Edited by Evangelos Tsotsas and Arun S. Mujumdar

Modern Drying Technology Modern Drying Technology Edited by E. Tsotsas and A. Mujumdar

Other Volumes

Volume 1: Computational Tools at Different Scales ISBN: 978-3-527-31556-7

Volume 2: Experimental Techniques ISBN: 978-3-527-31557-4

Volume 3: Product Quality and Formulation ISBN: 978-3-527-31558-1

Volume 4: Energy Savings ISBN: 978-3-527-31559-8

Modern Drying Technology Set (Volumes 1 – 5) ISBN: 978-3-527-31554-3 Edited by Evangelos Tsotsas and Arun S. Mujumdar

Modern Drying Technology

Volume 5: Process Intensification The Editors: All books published by Wiley-VCH are carefully produced. Nevertheless, authors, editors, and Prof. Evangelos Tsotsas publisher do not warrant the information contained Otto von Guericke University in these books, including this book, to be free of Thermal Process Engineering errors. Readers are advised to keep in mind that Universitätsplatz 2 statements, data, illustrations, procedural details or 39106 Magdeburg other items may inadvertently be inaccurate. Germany Library of Congress Card No.: applied for

Prof. Arun S. Mujumdar British Library Cataloguing-in-Publication Data National University of Singapore A catalogue record for this book is available from the Mechanical Engineering/Block EA 07-0 British Library. 9 Engineering Drive 1 Bibliographic information published by the Deutsche Singapore 117576 Nationalbibliothek Singapore The Deutsche Nationalbibliothek lists this publication in the Deutsche Nationalbibliografie; detailed bibliographic data are available on the Internet at < http:// dnb.d-nb.d e> .

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Contents

Series Preface XI Preface of Volume 5 XV List of Contributors XIX Recommended Notation XXIII EFCE Working Party on Drying: Address List XXIX

1 Impinging Jet Drying 1 Eckehard Specht 1.1 Application 1 1.2 Single Nozzle 4 1.3 Nozzle Fields 7 1.3.1 Arrays of Single Nozzles 7 1.3.2 Hole Channels 12 1.3.3 Perforated Plates 13 1.3.4 Nozzles for Cylindrical Bodies 14 1.4 Summary of the Nusselt Functions 16 1.5 Design of Nozzle Field 17 1.6 Conclusion 23 References 24

2 Pulse Combustion Drying 27 Ireneusz Zbicinski, Tadeusz Kudra, and Xiangdong Liu 2.1 Principle of Pulse Combustion 27 2.2 Pulse Combustors: Design and Operation 32 2.2.1 Pulse Combustors with Mechanical Valves 32 2.2.2 Pulse Combustors with Aerodynamic Valves 34 2.2.3 Frequency-Tunable Pulsed Combustors 35 2.3 Aerodynamics, Heat and Mass Transfer 36 2.3.1 Atomization 37 2.3.2 Heat and Mass Transfer 38 2.4 Modeling of Pulse Combustion Drying 42 2.5 Pulse Combustion in Drying 48 References 53 VIj Contents 3 Superheated Steam Drying of Foods and Biomaterials 57 Sakamon Devahastin and Arun S. Mujumdar 3.1 Introduction 57 3.2 Principle of Superheated Steam Drying (SSD) 58 3.3 Atmospheric-Pressure Superheated Steam Drying 61 3.4 Low-Pressure Superheated Steam Drying (LPSSD) 69 3.5 Application of LPSSD to Improve the Quality of Foods and Biomaterials 76 3.6 Concluding Remarks 82 References 83

4 Intensification of Fluidized-Bed Processes for Drying and Formulation 85 Evangelos Tsotsas, Stefan Heinrich, Michael Jacob, Mirko Peglow, and Lothar M€orl 4.1 Introduction 85 4.2 Intensification by Apparatus and Flow Design 86 4.2.1 Different Types of Spouted Bed 86 4.2.2 Operating Characteristics of Spouted Beds 93 4.2.3 Mass and Heat Transfer in ProCell Units 100 4.2.4 Discrete Particle Modeling 107 4.3 Intensification by Contact Heating 112 4.3.1 General Principle 112 4.3.2 Main Effects and Influences 114 4.3.3 Further Remarks on Modeling 121 4.4 Further Methods of Intensification 126 4.5 Conclusion 127 References 128

5 Intensification of Freeze-Drying for the Pharmaceutical and Food Industries 131 Roberto Pisano, Davide Fissore, and Antonello A. Barresi 5.1 Introduction 131 5.2 Exergetic Analysis (and Optimization) of the Freeze-Drying Process 133 5.3 Process Intensification in Vacuum Freeze-Drying of Liquids 139 5.3.1 Regulation of Nucleation Temperature During Freezing 140 5.3.2 Use of Organic Solvents and Cosolvents 144 5.4 Atmospheric Freeze-Drying 146 5.5 Use of Combined Technologies for Drying Heat-Sensitive Products 150 5.5.1 Microwave-Assisted Drying 150 5.5.2 Ultrasound-Assisted Drying 152 5.6 Continuous Freeze-Drying 154 5.7 Conclusions 155 References 157 Contents jVII 6 Drying of Foamed Materials 163 Ireneusz Zbicinski, Julia Rabaeva, and Artur Lewandowski 6.1 Introduction 163 6.2 Foam Properties 164 6.3 Foam Spray Drying 167 6.3.1 Processing Principles 167 6.3.2 Final Product Properties 172 6.4 Foam-Mat Drying 181 6.5 Summary 187 References 188

7 Process-Induced Minimization of Mass Transfer Barriers for Improved Drying 191 Henry J€ager, Katharina Sch€ossler, and Dietrich Knorr 7.1 Introduction 191 7.2 Structural Characterization of Plant Raw Materials and Impact of PEF and Ultrasound 192 7.2.1 Methods for Analysis of Tissue Structure and Quantification of Cell Damage 192 7.2.2 PEF: Principles and Impact on Plant Tissue Structure 195 7.2.2.1 Introduction to PEF Technology 195 7.2.2.2 PEF: Impact on Plant Tissue Structure 196 7.2.3 Ultrasound: Principles and Impact on Plant Tissue Structure 199 7.2.3.1 Introduction to Ultrasound Technology 199 7.2.3.2 Ultrasound: Impact on Plant Tissue Structure 200 7.3 Pulsed Electric Field (PEF) Application as a Pretreatment 204 7.3.1 Osmotic Dehydration 205 7.3.2 Air Drying 206 7.3.3 Impact of PEF on Freezing and Freeze-Drying Behavior of Raw Materials 208 7.3.4 Quality Characteristics Affected by PEF Pretreatment 211 7.4 Contact Ultrasound for Combined Drying Processes 216 7.4.1 Ultrasound in Osmotic Dehydration 217 7.4.2 Contact Ultrasound in Air Drying 218 7.4.3 Contact Ultrasound in Freeze-Drying 221 7.4.4 Quality Characteristics Affected by Ultrasound-Combined Drying Processes 224 7.5 Conclusion 226 References 230

8 Drying Assisted by Power Ultrasound 237 Juan Andres Carcel, Jose Vicente García-Perez, Enrique Riera, Carmen Rossello, and Antonio Mulet 8.1 Introduction 237 8.2 Ultrasound 239 VIIIj Contents 8.2.1 Ultrasound Waves 239 8.2.1.1 Power 239 8.2.1.2 Frequency 240 8.2.1.3 Attenuation 240 8.2.1.4 Acoustic Impedance 240 8.2.2 Effects of Ultrasound on Mass Transfer 241 8.3 Ultrasonic Equipment 242 8.3.1 Source of Energy 243 8.3.2 Transducers 243 8.3.3 Application Systems 245 8.3.3.1 Treatments in Liquid Media 245 8.3.3.2 Treatments in Gas Media 247 8.4 Influence of the Main Process Variables on Drying Intensification by Ultrasound 250 8.4.1 Ultrasonic Power Applied 250 8.4.1.1 Ultrasonic Field Measurements 251 8.4.1.2 Ultrasonic Intensity and Effects 252 8.4.1.3 Influence of the Characteristics of the Medium on Ultrasonic Intensity 258 8.4.2 Drying Air Temperature 263 8.4.3 Ultrasound–Sample Interaction 266 8.5 Conclusions 272 References 273

9 Microwave-Assisted Drying of Foods – Equipment, Process and Product Quality 279 Yingqiang Wang, Min Zhang, and Arun S. Mujumdar 9.1 Introduction 279 9.2 Microwave-Assisted Drying of Foods 281 9.2.1 Basic Principles of Microwave-Assisted Drying 281 9.2.2 Energy Absorption by Products During Dielectric Heating 283 9.2.3 Dielectric Properties 283 9.2.4 Penetration Depth 285 9.3 Microwave-Assisted Drying Equipment 285 9.3.1 Microwave-Assisted Convective Drying Equipment 286 9.3.2 Microwave-Assisted Vacuum Drying Equipment 287 9.3.3 Microwave-Assisted Freeze-Drying Equipment 290 9.3.4 Microwave-Assisted Spouted Bed Drying Equipment 291 9.4 Microwave-Assisted Drying Process 292 9.4.1 Moisture Loss 293 9.4.2 Temperature Distributions 295 9.4.2.1 Temperature Variations at Fixed Levels of Microwave Power 296 9.4.2.2 Temperature Variations at Variable Microwave Power without Controlling Temperature 298 Contents jIX 9.4.2.3 Temperature Change with Time-Adjusted Power in Feedback Temperature Control 299 9.4.3 Energy Consumption 299 9.4.4 Dielectric Breakdown 302 9.4.5 Changes in Dielectric Properties 304 9.4.6 Quality Changes in Food during Microwave-Assisted Drying 305 9.5 Microwave-Assisted Drying Process Control and Optimal Operation 308 9.5.1 Factors Controlling Microwave-Assisted Drying Processes 308 9.5.2 Optimal Operation Strategy 308 9.6 Concluding Remarks 310 References 312

10 Infrared Drying 317 German Efremov 10.1 Introduction 317 10.2 Radiation Heat Transfer 318 10.2.1 General Principles 318 10.2.2 Reflection, Absorption, and Transmission 319 10.2.3 Infrared Spectrum 321 10.3 Classification, Research, and Applications of Radiation Drying 323 10.3.1 Classification 323 10.3.2 Solar Drying 325 10.3.3 Infrared Drying 326 10.3.4 Catalytic Infrared Drying 329 10.4 Types of Radiators 332 10.4.1 General Considerations 332 10.4.2 Electric Radiators 333 10.4.3 Gas-Heated IR Radiators 335 10.5 Interaction between Matter and Infrared Radiation 337 10.5.1 General Relationships 337 10.5.2 Radiation Properties of Materials 339 10.6 Kinetics of Infrared Drying 342 10.7 Infrared Drying Combined with other Types of Drying 345 10.7.1 IR and Convective Drying 346 10.7.2 IR and Microwave Drying 347 10.7.3 IR and Freeze-Drying 348 10.7.4 IR with other Types of Drying 348 10.8 Conclusions 351 References 352

Index 357

jXI

Series Preface

The present series is dedicated to drying, that is, to the process of removing moisture from solids. Drying has been conducted empirically since the dawn of the human race. In traditional scientific terms it is a unit operation in chemical engineering. The reason for the continuing interest in drying and, hence, the motivation for the series, concerns the challenges and opportunities. A permanent challenge is connected to the sheer amount and value of products that must be dried – either to attain their functionalities, or because moisture would damage the material during subsequent processing and storage, or simply because customers are not willing to pay for water. This comprises almost every material used in solid form, from foods to pharmaceuticals, from minerals to detergents, from polymers to paper. Raw materials and commodities with a low price per kilogram, but with extremely high production rates, and also highly formulated, rather rare but very expensive specialties have to be dried. This permanent demand is accompanied by the challenge of sustainable devel- opment providing welfare, or at least a decent living standard, to a still-growing humanity. On the other hand, opportunities emerge for drying, as well as for any other aspect of science or living, from either the incremental or disruptive develop- ment of available tools. This duality is reflected in the structure of the book series, which is planned for five volumes in total, namely: Volume 1: Computational tools at different scales Volume 2: Experimental techniques Volume 3: Product quality and formulation Volume 4: Energy savings Volume 5: Process intensification. As the titles indicate, we start with the opportunities in terms of modern computational and experimental tools in Volumes 1 and 2, respectively. How these opportunities can be used in fulfilling the challenges, in creating better and new products, in reducing the consumption of energy, in significantly improving existing or introducing new processes will be discussed in Volumes 3, 4 and 5. In this sense, the first two volumes of the series will be driven by science; the last XIIj Series Preface three will try to show how engineering science and technology can be translated into progress. In total, the series is designed to have both common aspects with and essential differences from an extended textbook or a handbook. Textbooks and handbooks usually refer to well-established knowledge, prepared and organized either for learning or for application in practice, respectively. On the contrary, the ambition of the present series is to move at the frontier of “modern drying technology”, describing things that have recently emerged, mapping things that are about to emerge, and also anticipating some things that may or should emerge in the near future. Consequently, the series is much closer to research than textbooks or handbooks can be. On the other hand, it was never intended as an anthology of research papers or keynotes – this segment being well covered by periodicals and conference proceedings. Therefore, our continuing effort will be to stay as close as possible to a textbook in terms of understandable presentation and as close as possible to a handbook in terms of applicability. Another feature in common with an extended textbook or a handbook is the rather complete coverage of the topic by the entire series. Certainly, not every volume or chapter will be equally interesting for every reader, but we do hope that several chapters and volumes will be of value for graduate students, for researchers who are young in age or thinking, and for practitioners from industries that are manu- facturing or using drying equipment. We also hope that the readers and owners of the entire series will have a comprehensive access not to all, but to many significant recent advances in drying science and technology. Such readers will quickly realize that modern drying technology is quite interdisciplinary, profiting greatly from other branches of engineering and science. In the opposite direction, not only chemical engineers, but also people from food, mechanical, environ- mental or medical engineering, material science, applied chemistry or physics, computing and mathematics may find one or the other interesting and useful results or ideas in the series. The mentioned interdisciplinary approach implies that drying experts are keen to abandon the traditional chemical engineering concept of unit operations for the sake of a less rigid and more creative canon. However, they have difficulties of identifi- cation with just one of the two new major trends in chemical engineering, namely process-systems engineering or product engineering. Efficient drying can be completely valueless in a process system that is not efficiently tuned as a whole, while efficient processing is certainly valueless if it does not fulfill the demands of the market (the customer) regarding the properties of the product. There are few topics more appropriate in order to demonstrate the necessity of simultaneous treatment of product and process quality than drying. The series will try to work out chances that emerge from this crossroads position. One further objective is to motivate readers in putting together modules (chapters from different volumes) relevant to their interests, creating in this manner individ- ual, task-oriented threads trough the series. An example of one such thematic thread set by the editors refers to simultaneous particle formation and drying, with a focus Series Preface jXIII on spray fluidized beds. From the point of view of process-systems engineering, this is process integration – several “unit operations” take place in the same equipment. On the other hand, it is product engineering, creating structures – in many cases nanostructures – that correlate with the desired application properties. Such properties are distributed over the ensemble (population) of particles, so that it is necessary to discuss mathematical methods (population balances) and numerical tools able to resolve the respective distributions in one chapter of Volume 1. Measuring techniques providing access to properties and states of the particle system will be treated in one chapter of Volume 2. In Volume 3, we will attempt to combine the previously introduced theoretical and experimental tools with the goal of product design. Finally, important issues of energy consumption and process intensification will appear in chapters of Volumes 4 and 5. Our hope is that some thematic combinations we have not even thought about in our choice of contents will arise in a similar way. As the present series is a series of edited books, it can not be as uniform in either writing style or notation as good textbooks are. In the case of notation, a list of symbols has been developed and will be printed in the beginning of every volume. This list is not rigid but foresees options, at least partially accounting for the habits in different parts of the world. It has been recently adopted as a recommendation by the Working Party on Drying of the European Federation of Chemical Engineering (EFCE). However, the opportunity of placing short lists of additional or deviant symbols at the end of every chapter has been given to all authors. The symbols used are also explained in the text of every chapter, so that we do not expect any serious difficulties in reading and understanding. The above indicates that the clear priority in the edited series was not in uniformity of style, but in the quality of contents that are very close to current international research from academia and, where possible, also from industry. Not every potentially interesting topic is included in the series, and not every excellent researcher working on drying contributes to it. However, we are very confident about the excellence of all research groups that we were able to gather together, and we are very grateful for the good cooperation with all chapter authors. The quality of the series as a whole is set mainly by them; the success of the series will primarily be theirs. We would also like to express our acknowl- edgments to the team of Wiley-VCH who have done a great job in supporting the series from the first idea to realization. Furthermore, our thanks go to Mrs Nicolle Degen for her additional work, and to our families for their tolerance and continuing support. Last but not least, we are grateful to the members of the Working Party on Drying of the EFCE for various reasons. First, the idea about the series came up during the annual technical and business meeting of the working party 2005 in Paris. Secondly, many chapter authors could be recruited among its members. Finally, the Working Party continues to serve as a panel for discussion, checking and readjustment of our conceptions about the series. The list of the members of the working party with their affiliations is included in every volume of the series in the sense of acknowledgment, XIVj Series Preface but also in order to promote networking and to provide access to national working parties, groups and individuals. The present edited books are complementary to the regular activities of the EFCE Working Party on Drying, as they are also comple- mentary to various other regular activities of the international drying community, including well-known periodicals, handbooks, and the International Drying Symposia.

December 2006 Evangelos Tsotsas Arun S. Mujumdar jXV

Preface of Volume 5

Volume 5 of “Modern Drying Technology” is dedicated to “Process intensification”. This is a natural conclusion for the series. “Computational tools at different scales” and “Experimental techniques”, presented in Vol. 1 and Vol. 2, respectively, were used to discuss “Product quality and formulation” in Vol. 3, and “Energy savings” in Vol. 4. Now the goal is not as specific as in Vol. 4, but more general and quite ambitious; namely, to use as intensive drying processes as possible in order to reduce the drying time and hence the equipment size. Insights from all previous volumes of the series must be implemented and applied to this purpose, leading to the following ten chapters of Vol. 5:

Chapter 1: Impinging jet drying Chapter 2: Pulse combustion drying Chapter 3: Superheated steam drying of foods and biomaterials Chapter 4: Intensification of fluidized-bed processes for drying and formulation Chapter 5: Intensification of freeze-drying for the pharmaceutical and food industries Chapter 6: Drying of foamed materials Chapter 7: Process-induced minimization of mass transfer barriers for improved drying Chapter 8: Drying assisted by power ultrasound Chapter 9: Microwave-assisted drying of foods – Equipment, process and product quality Chapter 10: Infrared drying

Frequent mention of foods, biomaterials and pharmaceuticals in the list of contents shows that process intensification is not a stand-alone perspective, but a challenge which usually must be addressed under serious constraints set by the quality requirements of valuable products that might suffer damage during processing. On the other hand, the list of contents of this final volume is longer than the lists of previous volumes in this series, indicating that a large variety of approaches and methods can lead to process intensification in practice, and that the international drying community is continually and persistently working on their further devel- opment and implementation. XVIj Preface of Volume 5 External, gas-side heat and mass transfer resistances can seriously limit the rate of drying processes, but they can also be radically reduced by high velocity flow impingement on the surface of the material to be dried. This is easy to implement for flat products such as paper, textiles, tissues, tiles or wood veneer, but connected with questions about how many nozzles shall be used, and how these nozzles shall be placed and operated. Answers to those questions are provided in Chapter 1, in a concise way that refrains from the consideration of less significant details and aims at immediate engineering applicability. The same purpose of lowering external heat and mass transfer resistances can be achieved by imposing, instead of a steady turbulent flow, an oscillatory flow of drying gas around the material to be dried. This can be realized by drying in flue gas coming from a special, pulse burner via a tailpipe to the drying chamber. Intensification of the external heat and mass transfer maynotbeasspectacularasincaseof impinging jets, but the method is applicable to virtually any kind of convective dryer, i.e. it is not restricted to flat products. Construction and operation of the respective combustors, enhancement of heat and mass transfer, and modeling are discussed in Chapter 2. In Chapter 3, the focus is shifted to the use of superheated steam, instead of hot air of flue gas, as the drying agent, which has an influence on both, the external and the internal heat and mass transfer. A major advantage of this process is that energy can much easier be recovered from exhaust steam, than from the wet exhaust air of conventional drying processes. The necessity of operating above the boiling point of the liquid to be removed, usually water, may be turned into an advantage by combining the drying process with, e.g., sterilization or cooking of foods and biomaterials. Damage that such materials might suffer at the boiling temperature of water at ambient pressure can be prevented by reducing the operating pressure, i.e. by low-pressure superheated steam drying. Alternatively, drying rates can be boosted in fluidized beds by combining heat transfer from the fluidization air with indirect heat transfer, usually from immersed steam tubes. Fundamentals and applications of respective processes are discussed in Chapter 4. It is pointed out that the resulting process intensification can be used for increasing the capacity of the dryer, or for reducing the temperature level in order to protect thermally sensitive products. Moreover, the process can be significantly intensified by applying spouted beds, instead of conventional fluidized beds. The background of this behavior is that regions of extremely high gas velocity can be realized in specially designed spouted beds with adjustable air inlet. However, there are foods and pharmaceuticals which are so sensitive, that they must be dried from the frozen state. Purposeful use of the notoriously slow process of freeze drying increases the necessity and urgency of process intensification measures. Various such measures are available, as discussed in Chapter 5, including automatic control for better drying cycles, favorable templating of the solid matrix to be dried by controlled freezing, the use of organic solvents, freeze drying under atmospheric conditions, or the transition from batch to the continuous operation mode. Moreover, hybrid processes can be applied, such as microwave or ultrasound assisted freeze drying. Preface of Volume 5 jXVII Sometimes, product quality requirements meet with the goal of more intense drying processes for hard-to-dry products, such as fruit pulps, juices, or dairy. Foam drying techniques provide attractive solutions for such cases. Foamed products can be produced in spray dryers by injecting inert gas to the feed of the dryer, before or during atomization. Alternatively, solutions or dispersions can be whipped to foam that is subsequently dried in any appropriate type of equipment, which is denoted by foam-mat drying. Respective process configurations, enhancement of drying by increased surface area and more open structures, and resulting product properties are discussed in Chapter 6. In some other cases, biological materials to be dried contain natural barriers to mass transfer, such as cell membranes. Then, drying can be enhanced by applying pulsed electric fields to create pores in the membranes or disintegrate the cells, followed by osmotic dehydration, hot-air drying, or freeze drying. Similar effects can be attained by application of ultrasound to support and assist the mentioned drying processes. Principles and results of these novel technologies are presented in Chapter 7, along with methods for the structural and textural characterization of the materials, and quality characteristics of the resulting products. A more detailed treatment of the application of ultrasound is provided in Chapter 8, along with a discussion of the principles of generation and transmission of ultrasound energy to the material to be treated. It is pointed out that power ultrasound can be used to assist both, liquid-solid processes, such as brine treatment, and drying. The acoustic field is shown to enhance, by a number of mechanisms, both, the external and the internal mass transfer when combined with hot air or atmospheric freeze drying of vegetables and fruits. The more porous the material, and the lower the permissible temperature and gas velocity, the higher is the intensification that can be reached by application of power ultrasound. Another method of hybrid or assisted processing is to support hot-air drying, vacuum drying, freeze drying, or spouted bed drying by microwaves. Microwaves have the unique property of targeting heat supply to the consumer, i.e. to the wet interior of drying materials. Respective processes, equipment, and the enhancement of drying rate that can be achieved by means of the microwaves are thoroughly presented in Chapter 9. Moreover, issues of energy consumption, automatic control, and product quality are addressed. Agricultural products and food materials are, again, in the focus of the discussion. Infrared radiation usually does not penetrate deep into materials, but it can significantly intensify drying processes by supplying significant and well controlla- ble amounts of energy to the surface. In a comprehensive treatment of infrared drying in Chapter 10 different types of radiators, including gas-fired ones, are presented, the necessity of matching the infrared spectrum used with the properties of the material to be dried is stressed, and opportunities to improve product quality by intermittent radiation supply are pointed out. Combinations of infrared heat supply with hot-air drying, microwave drying, freeze drying, and heat-pump drying are discussed. Volume 5 brings several thematic threads set in previous volumes, for instance on fluidized bed drying or food processing, to their contemporary completion. It XVIIIj Preface of Volume 5 reflects the interdisciplinary and multi-scale character of modern drying technology in a similar way as the previous volumes of the series. Therefore, we hope that this final volume and the entire series have at least partially attained the goal of providing all people working on drying in industry and research with a map that shows where drying science and technology are, where they are presently growing to cope with increasing challenges and application demands, and where they may be in the future. In other words, we hope that the series can contribute to the solution of specific, well defined practical tasks (“improve quality, save energy, cut costs”,as promised in the flyer of the publisher), but that it can also motivate further exploratory work and inspire to innovation in an important and rewarding field of engineering science. In this farewell preface, we would like to renew our profound acknowledgement of all persons who have made this series possible - our families and co-workers, the excellent editorial team of the publisher, all our outstanding and esteemed colleagues and friends who have served as chapter authors, the countless engineers and scientists whose contributions are quoted in the book series, but also those who have contributed to drying science and practice without finding individual citation.

Summer 2013 Evangelos Tsotsas Arun S. Mujumdar jXIX

List of Contributo rs

Editors Prof. Juan Andes r arcel C Universidad Polecnicait de Valencia Prof. Evangelos Tsotsas Departamento de Tecnología de Otto von Guericke University Alimentos Magdeburg Cami de Vera s/n Thermal Process Engineering 46022 Valencia PSF 4120 Spain 39106 Magdeburg Email: [email protected] Germany Prof. Sakamon Devahastin Email: [email protected] King Mongkut’ s University of Technology Thonburi Prof. Arun S. Mujumdar Department of Food Engineering McGill University 126 Pracha u-tid Road Department of Bioresource Bangkok 10140 Engineering Thailand 2111 Lakeshore Road Email: [email protected] Sainte-Anne-de-Bellevue Quebec H9X 3V9 Prof. German Efremov Canada Moscow State Open University Email: [email protected] Street Krasnogo Mayaka 13a cor. 2, app. 71 Authors 117570 Moscow Russia Prof. Antonello A. Barresi Email: [email protected] Politecnico di Torino Dipartimento di Scienza Applicata e Dr. Davide Fissore Tecnologia Politecnico di Torino Corso Duca degli Abruzzi 24 Dipartimento di Scienza Applicata e 10129 Torino Tecnologia Italy Corso Duca degli Abruzzi 24 Email: [email protected] 10129 Torino Italy Email: davide. fi[email protected] XXj List of Contributors Prof. Joe s Vicente García-erezP Dr. Tadeusz Kudra Universidad Poliecnicat de Valencia CanmetENERGY Departamento de Tecnología de 957 de Salieres Alimentos St. Jean-sur-Richelieu Cami de Vera s/n Quebec J2W 1A3 46022 Valencia Canada Spain Email: [email protected] Email: [email protected] Artur Lewandowski Prof. Stefan Heinrich Lodz University of Technology Hamburg University of Technology Faculty of Process and Environmental Institute of Solids Process Engineering Engineering and Particle Technology ul. Wolczanska 213 Denickestrasse 15 93-924 Lodz 21073 Hamburg Poland Germany Email: [email protected] Email: [email protected] Prof. Xiangdong Liu Dr. Michael Jacob China Agricultural University Glatt Ingenieurtechnik GmbH College of Engineering Nordstrasse12 17 Qinghua East Rd. 99427 Weimar Beijing 100083 Germany P. R. China Email: [email protected] Email: [email protected]

Dr. Henryager€ J Prof. Lothar €orlM Technische Universi€att Berlin Otto von Guericke University Department of Food Biotechnology Magdeburg and Food Process Engineering Chemical Equipment Design Koenigin-Luise-Strasse 22 PSF 4120 14195 Berlin 39106 Magdeburg Germany Germany Email: [email protected] Email: [email protected]

Prof. Dietrich Knorr Prof. Arun S. Mujumdar Technische Universi€att Berlin McGill University Department of Food Biotechnology Department of Bioresource and Food Process Engineering Engineering Koenigin-Luise-Strasse 22 2111 Lakeshore Road 14195 Berlin Sainte-Anne-de-Bellevue Germany Quebec H9X 3V9 Email: [email protected] Canada Email: [email protected] List of Contributors jXXI Prof. Mirko Peglow Prof. Carmen Rossello IPT-PERGANDE GmbH English University of Illes Balears Wilfried-Pergande-Platz 1 Spanish Departamento 06369 Weißandt-Golzau€ de Química Germany Ctra. Valldemossa km 7.5 Email: [email protected] 07122 Palma Mallorca Spain Dr. Roberto Pisano Email: [email protected] Politecnico di Torino Dipartimento di Scienza Applicata e Dr. Katharina €osslerSch Tecnologia Technische Universi€att Berlin Corso Duca degli Abruzzi 24 Department of Food Biotechnology 10129 Torino and Food Process Engineering Italy Koenigin-Luise-Strasse 22 Email: [email protected] 14195 Berlin Germany Prof. Antonio Mulet Pons Email: [email protected] Universidad Poliecnicat de Valencia Departamento de Tecnología de Alimentos Prof. Eckehard Specht Cami de Vera s/n Otto von Guericke University 46022 Valencia Magdeburg Spain Thermodynamics and Combustion Email: [email protected] PSF 4120 39106 Magdeburg Dr. Julia Rabaeva Germany Lodz University of Technology Email: [email protected] Faculty of Process and Environmental Engineering Prof. Evangelos Tsotsas ul. Wolczanska 213 Otto von Guericke University 93-924 Lodz Magdeburg Poland Thermal Process Engineering Email: [email protected] PSF 4120 39106 Magdeburg Dr. Enrique Riera Germany Instituto de Seguridad de la Email: [email protected] Informaci on (ISI) Grupo de Sistemas Ultrasonicos CSIC Serrano 144 28006 Madrid Spain Email: [email protected] XXIIj List of Contributors Dr. Yingqiang Wang Prof. Ireneusz Zbicinski Jiangnan University Lodz University of Technology School of Food Science and Faculty of Process and Environmental Technology Engineering 214122 Wuxi ul. Wolczanska 213 Jiangsu Province 93-924 Lodz P. R. China Poland Email: [email protected] and Prof. Min Zhang Longdong University Jiangnan University College of Agriculture and Forestry School of Food Science and Lanzhou Road 45 Technology 745000 Qingyang 214122 Wuxi Gansu Province Jiangsu Province P. R. China P. R. China Email: [email protected] Email: [email protected] jXXIII

Recommended Notation

Alternative symbols are given in brackets Vectors are denoted by bold symbols, a single bar, an arrow or an index (e.g., index: i) Tensors are denoted by bold symbols, a double bar or a double index (e.g., index: i, j) r Multiple subscripts should be separated by colon (e.g., p;dry: density of dry particle)

A surface area m 2 a – w water activity B nucleation rate kg 1 m 1 s 1 b breakage function m 3 C (K) constant or coefficient various c specific heat capacity J kg 1K 1 D equipment diameter m D (d) diffusion coefficient m2 s 1 d diameter or size of solids m E energy J F mass flux function – FðV_ Þ volumetric flow rate m 3s 1 f relative (normalized) drying rate – f multidimensional number density – G shear function or modulus Pa G growth rate kg s 1 g acceleration due to gravity m s 2 H height m H enthalpy J H Heaviside step function – h specific enthalpy (dry basis) J kg 1 h (a) heat-transfer coefficient W m 2 K 1 ~ 1 h ðhNÞ molar enthalpy J mol 1 Dhv specific enthalpy of evaporation J kg I total number of intervals – XXIVj Recommended Notation J numerical flux function – J Jacobian matrix various j ðm_ ; JÞ mass flux, drying rate kg m 2 s 1 K dilatation function or bulk modulus Pa k (b) mass transfer coefficient m s 1 L length m M (m) mass kg M~ M M 1 ( , N) molecular mass kg kmol M_ ðWÞ mass flow rate kg s 1 m_ ðJ; jÞ mass flux, drying rate kg m 2 s 1 m_ volumetric rate of evaporation kg m 3 s 1 N number – N molar amount mol _ 1 N ðWN Þ molar flow rate mol s n molar density, molar concentration mol m 3 n number density m 3 n outward normal unit vector 2 1 n_ ðJN Þ molar flux mol m s P power W P total pressure kg m 1 s 2 p partial pressure/vapor pressure kg m 1 s 2 of component Q_ ðQÞ heat flow rate W q_ ðqÞ heat flux W m 2 R equipment radius m R individual gas constant J kg 1 K 1 ~ 1 1 R ðRN Þ universal gas constant J kmol K r radial coordinate m r pore (throat) radius m S saturation – S selection function s 1 s boundary-layer thickness m T temperature K,C t time s u velocity, usually in z-direction m s 1 u displacement m V volume, averaging volume m3 V_ ðFÞ volumetric flow rate m3 s 1 v specific volume m3 kg 1 v general velocity, velocity in x-direction m s 1 W weight force N W ðM_ Þ mass flow rate kg s 1 w velocity, usually in y-direction m s 1 X solids moisture content (dry basis) – x mass fraction in liquid phase – Recommended Notation jXXV x particle volume in population balances m3 x general Eulerian coordinate, coordinate m (usually lateral) x 0 general Lagrangian coordinate m ~ x ðxNÞ molar fraction in liquid phase – Y gas moisture content (dry basis) – y spatial coordinate (usually lateral) m y (v) mass fraction in gas phase – ~y ðy Þ – N molar fraction in gas phase z spatial coordinate (usually axial) m Operators ! gradient operator ! divergence operator D difference operator Greek letters a (h) heat-transfer coefficient W m 2 K 1 b (k) mass-transfer coefficient m s 1 b aggregation kernel s 1 d Dirac-delta distribution d (D) diffusion coefficient m2 s 1 e voidage – e emissivity – e small-scale parameter for periodic media – e strain – h efficiency – u angle, angular coordinate rad k thermal diffusivity m2 s 1 l thermal conductivity W m 1K 1 m dynamic viscosity kg m 1 s 1 m moment of the particle-size distribution various n kinematic viscosity m2 s 1 p circular constant – r density, mass concentration kg m 3 S summation operator s surface tension N m 1 s Stefan–Boltzmann constant for radiative Wm 2 K 4 heat transfer s standard deviation (of pore-size distribution) m s stress Pa t dimensionless time – F characteristic moisture content – w relative humidity – w phase potential Pa v angular velocity rad v (y) mass fraction in gas phase – XXVIj Recommended Notation Subscripts a at ambient conditions as at adiabatic saturation conditions b bound water bed bed c cross section c capillary cr at critical moisture content D drag dry dry dp at dewpoint eff effective eq equilibrium (moisture content) f friction g gas (dry) H wet (humid) gas I inner i, 1, 2, . . . component index, particle index i, j, k coordinate index, i,j,k = 1 to 3 in inlet value l liquid (alternative: as a superscript) m mean value max maximum mf at minimum fluidization min minimum N molar quantity o outer out outlet value P at constant pressure p particle pbe population balance equation ph at the interface r radiation rel relative velocity s solid (compact solid phase), alternative: as a superscript S at saturation conditions surf surface V based on volume v vapor, evaporation w water w wall wb at wet-bulb conditions wet wet 1 at large distance from interface Recommended Notation jXXVII Superscripts, Special symbols v volumetric strain rheological strain at saturation conditions or hi average, phase average a a or hi intrinsic phase average spatial deviation variable

jXXIX

EFCE Working Party on Drying: Add ress List

Prof. Odilio Alves-Filho Prof. Antonello Barresi (delegate) Norwegian University of Science and Dip. Scienza dei Materiali e Technology Ingegneria Chimica Department of Energy and Process Politecnico di Torino Engineering Corso Duca degli Abruzzi 24 Kolbjørn Hejes vei 1B 10129 Torino 7491 Trondheim Italy Norway [email protected] odilo. fi[email protected] Dr. Rainer Bellinghausen (delegate) Prof. Julien Andrieu (delegate) Bayer Technology Services GmbH UCB Lyon I/ESCPE BTS-PT-PT-PDSP LAGEP UMR CNRS 5007 batiment Building E 41 308 G 51368 Leverkusen 43 boulevard du 11 novembre 1918Germany 69622 Villeurbanne cedex [email protected] France [email protected] Dr. Carl-Gustav Berg Dr. Paul Avontuur (guest industry) Abo Akademi Glaxo Smith Kline Process Design Laboratory New Frontiers Science Park H89 Biskopsgatan 8 Harlow CM19 5AW 20500 Abo United Kingdom Finland fi [email protected] cberg@abo.

Prof. Christopher G. J. Baker Dr. Catherine Bonazzi (delegate) Drying Associates AgroParisTech - INRA Harwell International JRU for Food Process Engineering Business Centre 1 Avenue des Olympiades 404/13 Harwell Didcot 91744 Massy cedex Oxfordshire OX11 ORA France United Kingdom [email protected] [email protected] XXXj EFCE Working Party on Drying: Address List Paul Deckers M.Sc. (delegate) Prof. Dr. Istvan Farkas (chairman) Bodec Szent Istvan University Process Optimization and Dep. of Physics and Process Control Development Pater K. u. 1 Industrial Area ‘t Zand 2103 Godollo Bedrijfsweg 1 Hungary 5683 CM Best [email protected] The Netherlands [email protected] Dr. Dietrich Gehrmann Wilhelm-Hastrich-Str. 12 Henk van Deventer M.Sc. (delegate) 51381 Leverkusen TNO Germany Utrechtseweg 48 [email protected] 3704 HE Zeist The Netherlands Prof. Adrian-Gabriel Ghiaus (delegate) [email protected] Thermal Engineering Department Technical University of Civil Dr. German I. Efremov Engineering Pavla Korchagina 22 Bd. P. Protopopescu 66 129278 Moscow 021414 Bucharest Russia Romania [email protected] [email protected]

Prof. Trygve Eikevik Prof. Gheorghita Jinescu Norwegian University of Science University “Politehnica ” din and Technology Bucuresti Dep. of Energy and Process Department of Chemical Engineering Engineering Kolbjørn Hejes vei 1B 1 Polizu street, Room F210 7491 Trondheim 78126 Bucharest Norway Romania [email protected] [email protected]

Dr.-Ing. Ioannis Evripidis Prof. Dr. Gligor Kanevce Dow Deutschland GmbH & Co. OHGSt. Kliment Ohridski University P.O. Box 1120 Faculty of Technical Sciences 21677 Stade ul. Ivo Ribar Lola b.b. Germany Bitola [email protected] FYR of Macedonia [email protected] EFCE Working Party on Drying: Address List jXXXI Ir. Ian C. Kemp (delegate) Prof. Avi Levy (delegate) Glaxo SmithKline, R&D Ben-Gurion University of the Negev Gunnels Wood Road Department of Mechnical Stevenage SG1 2NY Engineering United Kingdom Beer-Sheva 84105 [email protected] Israel [email protected] Prof. Matthias Kind Karlsruhe Institute of Technology Jean-Claude Masson Institut ur€ f Thermische Rhodia Oerations,p CRTL-GI-DIPH Verfahrenstechnik 85 Avenue deseres F r Perret Kaiserstr. 12 69196 Saint-Font Cedex 76128 Karlsruhe France Germany [email protected] [email protected] Prof. Natalia Menshutina Prof. Stefan J. Kowalski Mendeleyev University of Chemical Poznan University of Technology Technology of Russia (MUCTR) Institute of Technology and Chemical High Technology Department Engineering Muisskaya sq.9 ul. Marii Sklodowskiej Cuvie 2 125047 Moscow 60965 Poznan Russia Poland [email protected] [email protected] Dr. Thomas Metzger Prof. Magdalini Krokida (delegate) BASF SE National Technical University of GCP/T T-L540 Athens 67056 Ludwigshafen Department of Chemical Germany Engineering [email protected] Politechnioupoli, Zografou Campus 15780 Athens Prof. Antonio Mulet Pons (delegate) Greece Universitat Politecnica de Valencia [email protected] Departament de Tecnologia d’Aliments Dr. Ir. Aneliqueg eonardL (delegate) Cami de Vera s/n UniversitedeLiege 46071 Valencia Departement de Chimie Appliqee u Spain Laboratoire de enieG Chimique [email protected] B^atiment B6c – Sart-Tilman 4000 Legei Belgium [email protected] XXXIIj EFCE Working Party on Drying: Address List Prof. Zdzislaw Pakowski (delegate) Prof. Carmen Rossello (delegate) Lodz University of Technology University of Illes Baleares Faculty of Process and Environmental Dep. Quimica Engineering Ctra. Valldemossa km 7.5 ul. Wolczanska 213 07122 Palma Mallorca 90924 Lodz Spain Poland [email protected] [email protected] Dr. Panayiotis Scarlatos Prof. Patrick Pere r(delegate) SusTchem Engineering LTD Ecole Centrale Paris 144 3rd September Street Laboratoire de enieG des Proedesc et 11251 Athens Ma eriauxt Greece Grande Voie des Vignes [email protected] 92295 C^atenay-h Malabry Dr. Michael onherr€ Sch France BASF SE [email protected] GCT/T – L 540 67056 Ludwigshafen Dr. Romain emondR Germany AgroParisTech – ENGREF [email protected] 14, Rue Girardet 54042 Nancy Dr. Andreas Schreiner France Novartis Pharma AG [email protected] WSJ-145.1.54 Lichtstr. 35 Dr. Roger Rent€omr 4056 Basel Karlstad University Switzerland Department of Environmental and [email protected] Energy Systems Universtietsgatan 2 Dr. Milan Stakic 65188 Karlstad Vinca9 Institute for Nuclear Sciences Sweden Center NTI [email protected] P.O. Box 522 11001 Belgrade Prof. Michel Roques Serbia Universite de Pau at des Pays [email protected] de l’ Adour 5 Rue Jules-Ferry Dr. Andrew Stapley (delegate) ENSGTI Loughborough University 64000 Pau Department of Chemical France Engineering [email protected] Loughborough, Leicestershire LE11 3TU United Kingdom [email protected] EFCE Working Party on Drying: Address List jXXXIII Prof. Stig Stenstrom (delegate) Prof. Evangelos Tsotsas (delegate) Lund University Otto von Guericke University Institute of Technology Thermal Process Engineering Department of Chemical P.O. Box 4120 Engineering 39016 Magdeburg P.O. Box 124 Germany 22100 Lund [email protected] Sweden [email protected] Thorvald Ullum M.Sc. GEA, Niro Prof. Ingvald Strommen (delegate) Gladsaxevej 305 Norwegian University of Science and 2860 Soeborg Technology Denmark Department of Energy and Process [email protected] Engineering Kolbjørn Hejes vei 1b Dr. Bertrand Woinet (delegate) 7491 Trondheim SANOFI-PASTEUR, CDP Norway B^atiment 8600 [email protected] 31-33 Quai Armand Barbes 69683 Neuville sur Saone^ cedex Prof. Czeslaw Strumillo (delegate) France Lodz University of Technology Bertrand.woinet@sanofi-aventis. Faculty of Process and Environmental com Engineering ul. Wolczanska 213 Prof. Ireneusz Zbicinski 90924 Lodz Lodz University of Technology Poland Faculty of Process and Environmental [email protected] Engineering ul. Wolczanska 213 Prof. Radivoje Topic (delegate) 90924 Lodz University of Belgrade Poland Faculty of Mechanical Engineering [email protected] 27 Marta 80 11000 Beograd Serbia [email protected] 1

1 Impinging Jet Drying Eckehard Specht

1.1 Application

Flat products such as tiles, tissue, paper, textiles and wood veneer are often dried using nozzle arrays (Mujumdar, 2007). Figure 1.1 illustrates the basic principle of the drying process using a nozzle array. Ambient air of temperature Ta is heated in a combustion chamber or in a heat exchanger to temperature T0, requiring the _ energy H0. The heated air is blown through the nozzle array in order to dry the product. These conditions result in the product temperature TS. The evaporating _ mass flow has the enthalpy Hv. The nozzle array is characterized by the inner diameter d of the individual nozzle, by the pitch t between the nozzles, and the velocity w of the released gas. There are basically three possible designs of nozzle arrays which differ with regard to the spent flow of the air (Fig. 1.2). In a field of individual nozzles the air can flow unimpeded between almost all nozzles; however, in a hole channel the air can flow only between those above. In a perforated plate the air can only continue to flow laterally and then escape. Hole channels and perforated plates are easier to produce than single nozzles, as they only require holes to be perforated. However, the heat transfer is the highest for nozzle fields and the lowest for perforated plates, as will be subsequently shown. For the design of the nozzle array the energy consumption needed for drying is essential. This is the energy for heating the air: _ ¼ r _ ðÞ ; ð : Þ H0 0V0cp T0 Ta 1 1 where T0 and Ta represent the temperatures of the heated air and the environment, fi respectively, and cP is the average speci c heat capacity between these two temperatures. The density and the volume flow rate refer to the heated air. Volume flow rate depends on the discharge velocity w and the number n of the nozzles with a diameter of d: p V_ ¼ n d2w: ð1:2Þ 0 4

Modern Drying Technology Volume 5: Process Intensification, First Edition. Edited by Evangelos Tsotsas and Arun S. Mujumdar. Ó 2014 Wiley-VCH Verlag GmbH & Co. KGaA. Published 2014 by Wiley-VCH Verlag GmbH & Co. KGaA. 2 1 Impinging Jet Drying

Fig. 1.1 Drying using a nozzle array.

The number of nozzles will depend on the nozzle pitch t and the surface area A of the material: A n ¼ : ð1:3Þ t2 The air temperature is calculated from the condition that the transferred heat has to cover the enthalpy of vaporization and the enthalpy to heat the dry material flow

from ambient temperature to the saturation temperature TS: _ _ a ATðÞ¼0 TS mvA D hv þ MscsðÞTS Ta : ð1:4Þ _ fl D In Eq. 1.4, mv is the evaporation ux, hv is the evaporation enthalpy, and cs is the specific heat capacity of the material. The evaporating mass flux is obtained from the relationship for the mass transfer _ P P Pa mv ¼ b ln : ð1:5Þ Rv TS P PS Here, the influence of one-side diffusion is taken into consideration. The gas

constant of the vapor is represented by Rv, the total pressure by p, the partial pressure of the vapor in the ambient air by pa, and the saturated vapor pressure by pS. Additionally, the analogy of the Nusselt and Sherwood function is applied to the ratio of the heat and mass transfer coefficients:

0:6 a ¼ brcPLe ; ð1:6Þ wherein for the exponent of the Prandtl number in the Nusselt function the value 0.4 was used. The saturation pressure is approximated from the equilibrium relationship Dhv 1 1 PS ¼ P0 exp ð1:7Þ Rv TS T0 1.1 Application 3

Fig. 1.2 Types of nozzle arrays. (a) Single-nozzle array; (b) Hole channel; (c) Perforated plate.

¼ ¼ with the reference condition P0,T0, for example, P0 1 bar, T0 373 K. The minimum required energy is the enthalpy of vaporization of the water _ _ Hv ¼ mv D hv: ð1:8Þ 4 1 Impinging Jet Drying _ The specific drying energy is the energy for heating the air H0 related to the enthalpy of evaporation. From the above equations results m_ D h pd2r c w v v þ T T H_ 0 P a S a pd2r c w 1 T T 0 ¼ ¼ 0 P þ S a : ð1:9Þ _ 2 _ D 2 a _ D Hv 4t mv hv 4t mv hv Theenthalpytoheatupthedrymaterialwasomittedforclaritypurposes.Specific drying energy consumption according to Eq. 1.9 is dependent on the heat transfer coefficient. Particularly at high rates of evaporation, the specific energy consumption is lower with a high heat transfer coefficient. This strongly influences the drying rate and the size of the apparatus and, as a result, an increasing heat transfer coefficient increases the rate of drying which in turn allows a reduction in the size of the apparatus. The setting and the regulation of the heat transfer coefficient is therefore of great importance. The heat transfer of bodies in a crossflow is relatively low. In generating a high heat transfer, nozzle arrays are implemented wherein the jet emerging from the nozzles is perpendicular to the body. Such flows are called stagnation point flows. Nozzles may be either round or slot-shaped. The fields of nozzles can be made from single nozzles, or hole channels, or from perforated plates with aligned or staggered arrangements, permitting a variety of geometric parameters. The heat transfer coefficient of nozzle arrays is therefore considered in more detail in the following.

1.2 Single Nozzle

First, an air jet emerging from a single nozzle is considered. In Fig. 1.3, the generated flow field of a nozzle is shown schematically. From the nozzle with the diameter d, the flow exits with the approximately constant speed w. The jet impinges the surface virtually unchanged with a constant velocity as long as

Fig. 1.3 Boundary layer of a stagnation point flow. 1.2 Single Nozzle 5

Fig. 1.4 Local heat transfer coefficients for a single nozzle.

therelativedistancebetweenthenozzleandthesurfaceh/dislessthan6.At greater distances the core velocity decreases reducing the heat transfer. In the impinged zone, also called the stagnation point region, the flow is redirected in a radial direction, and in this region the flow can always be considered as laminar. At the outer edge of the emerging jet, circular vortices form. In the deflection region at r/d 1, a large vortex forms but this disintegrates at approximately r/d 2 into many small vortices. Eventually, the flow becomes increasingly turbulent (refer to Angioletti et al., 2003, for photographs of the visualized flow). Increasing the nozzle distance from the stagnation point decreases the velocity, which in turn results in a decrease of the heat transfer coefficient. A completely laminar flow only occurs at Reynolds numbers smaller than 50; however, such small Reynolds numbers possess no meaning for technical heat-transfer processes. The local heat transfer coefficient depends on three geometric dimensions: the nozzle diameter d; the nozzle distance h; and the radial distance r. Figure 1.4 shows the local heat transfer coefficients as a function of relative radius r/d with the relative distance h/d as a parameter exemplarily for a nozzle with a diameter of 10 mm and an outlet velocity of 81 m s 1. It is evident for small distances of h/d ¼ 1 to 4 that the heat transfer coefficient at the stagnation point (r/d 0.5) is approximately constant, but then drops slightly to a relative minimum at r/d 1. In the region of the vortex degradation at r/d 2 the coefficient passes through a maximum and then decreases continuously. At radial distances r/d > 3.5, all profiles coincide for h/d 6. Only higher distances result in a lower heat transfer. For technical processes, the average heat transfer coefficient for a circular area,

ðr 1 a ¼ aðÞr 2prdr; ð1:10Þ pr2 0 6 1 Impinging Jet Drying

Fig. 1.5 Average Nusselt number of a single nozzle.

is of interest rather than the local heat transfer coefficient. The nozzle diameter exerts the greatest influence of all three geometric parameters. Therefore, the Nusselt number (Nu) and Reynolds number (Re) are defined as a ¼ d ; ¼ wd : ð : Þ Nu l Re n 1 11 In Fig. 1.5, the average Nusselt numbers are shown depending on the specific radius r/d, with the relative distance h/d as a parameter for a given Reynolds number. In the figure it is evident that from a distance of approximately r=d > 3, all profiles coincide for h=d 6; greater distances will then result in a lower heat transfer. As a consequence of Figs 1.4 and 1.5, it can be concluded that:

The relative nozzle distance should be h/d < 6 because for higher distances the heat transfer decreases. At a distance less than h/d ¼ 4, the local variation in the stagnation region increases. Up to relative stagnation point distances of approximately r/d ¼ 3, the average heat transfer is largely independent of the stagnation point distance. Only for greater distances will the heat transfer begin to decrease.

The Nusselt functions specified in the literature deviate from each other particularly with regard to the exponent of the Reynolds number; this is because the exponent depends on the radial distance. At the stagnation point the flow is always laminar, and the exponent is therefore 0.5. Turbulence fully develops only at a considerable distance from the stagnation point, and as a result the exponent must be 0.8. Figure 1.6 shows the increase of the exponent related to the radius (Adler, 2004). At a distance of r/d ¼ 3, the exponent has reached a value of only about 0.62. Different exponents are obtained depending on the size of the area considered in averaging. 1.3 Nozzle Fields 7

Fig. 1.6 Dependence of the exponent of the Reynolds number on distance.

The Nusselt functions reported in the literature for the stagnation point (r/d 0.5) can be approximated by

0:5 0:4 NuSt ¼ 0:72 Re Pr : ð1:12Þ For distances around r/d ¼ 3, the average heat transfer can be accurately approximated with a power of 0.67

: : Nu ¼ 0:12 Re0 67 Pr0 4: ð1:13Þ In the literature, values ranging from 0.67 to 0.7 are mostly reported and, in order to improve comparability, in all of the following discussions the Reynolds exponent is approximated by 0.67. Nusselt functions are typically determined for air, which is the typical medium used during application.

1.3 Nozzle Fields

Inthecaseoflargerareastobecooledorheated, several nozzles in a so-called nozzle fieldmustbeused.However,therelativenozzlepitcht/dmustnowbetaken into consideration as an additional geometric parameter. Initially, arrays of single nozzles are discussed, and it is assumed that the pitch is equal in both directions.

1.3.1 Arrays of Single Nozzles

fi The pro les of the local Nusselt number Nulo for a row of three nozzles is shown in Fig. 1.7 (Attalla, 2005). Placement of the nozzles is indicated in the figure, where it can be seen how the typical profiles of the single nozzles overlap. 8 1 Impinging Jet Drying

Fig. 1.7 Profiles of the local Nusselt number in an array of individual nozzles.

In Fig. 1.2a, the view of a nozzle array is shown with the impacted area. The nozzles may be arranged either in-line or staggered (in this figure the nozzles are in-line). In any case, each nozzle is influenced by the square area t2 of the nozzle pitch, and by using this area the average of the heat transfer can be determined. In Fig. 1.8, the average Nusselt number is shown in exemplary fashion as a function of the pitch t/d in an in-line assembly, with the relative nozzle distance h/ d as a parameter for two Reynolds numbers. It is evident that a pitch of t/d ¼ 6 will always result in a pronounced maximum. In Fig. 1.9, the average Nusselt number is shown as a function of the relative nozzle distance with the pitch as a parameter. The profiles for both the in-line and staggered arrangements are shown. The Nusselt numbers are approximately constant up to a relative nozzle distance of about five, and then decrease continuously. When considering the profiles of the average Nusselt number, no significant difference can be seen between the in-line and staggered arrangements.

Fig. 1.8 Average heat transfer dependent on the pitch of the individual nozzles. 1.3 Nozzle Fields 9

Fig. 1.9 Average heat transfer dependent on the relative nozzle distance.

However, the profiles of the local Nusselt number differ considerably (this point will be discussed further). For t/d ¼ 6 (maximum heat transfer), the Nusselt function in the stagnation area can be expressed as r/d 0.5

0:5 0:4 NuSt ¼ 0:82 Re Pr ; ð1:14Þ The average Nusselt number over the entire surface area for h/d < 5is

: : Nu ¼ 0:16 Re0 67 Pr0 4 ð1:15Þ (Attalla and Specht, 2009). In comparison with respect to the Nusselt functions for single nozzle (Eqs 1.12 and 1.13), it is evident that the heat transfer in the nozzle array is higher than that of the single nozzle under the condition of similar Reynolds numbers. Regarding the average heat transfer, this increase is approximately 30%. In the following, the influence of the nozzle arrangement is considered in more detail. In Fig. 1.10 infrared images of the temperature field of a metal plate for both nozzle arrangements are shown. The average heat transfer is equal in both arrangements, as long as the relative nozzle distance is equal, as shown previously. However, there are differences in the profiles of the local heat transfer coefficient, and thus in the uniformity of heat transfer. In both images two lines are shown where the heat transfer is either maximal or minimal. The maximum heat transfer always occurs along the line passing through the stagnation points of the nozzles; accordingly, the minimum heat transfer always occurs on the line passing the middle between the stagnation points. As a consequence, the heat transfer is different in the longitudinal and transverse directions on a belt that passes beneath the nozzle array. In Fig. 1.11 for a specific example, the profiles of the Nusselt number along the lines of the maximum and minimum heat transfer for the in-line arrangement and the staggered arrangement are shown, respectively. For the in-line arrangement, 10 1 Impinging Jet Drying

Fig. 1.10 Temperature field of a plate with (a) an in-line and (b) a staggered arrangement of the nozzles.

heat transfer along line A is always higher than along line B, so that the heat transfer is different in the cross-section. The high difference along line A is especially noteworthy. In this example the Nusselt number fluctuates between the values 19 and 3, which is approximately a factor of 6. Also on line B the Nusselt number fluctuates by about a factor of 4, and the heat transfer is therefore very uneven. In the staggered arrangement, both lines intersect so that the heat transfer takes alternatingly high and low values. It is also worth noting, that the differences

Fig. 1.11 Local heat transfer along lines A and B in (a) an in-line and (b) a staggered nozzle arrangement. 1.3 Nozzle Fields 11

Fig.1.12 Average heat transfer along lines A and B (cf. Fig. 1.10) in (a) an in-line and (b) a staggered arrangement.

between the maximum and minimum values of a line are significantly lower than for the in-line arrangement. On line A, the factor between the highest Nusselt number (95) and the lowest Nusselt number (32) amounted to only approximately three, compared to six for the in-line arrangement. Also, the difference between the maximum and minimum Nusselt numbers on line B for the staggered arrangement was less by a factor of 2 than for the in-line arrangement. In Fig. 1.12, along lines A and B, the average Nusselt numbers for both arrangements are shown as a function of the relative distance for two Reynolds numbers at the optimum nozzle pitch. Up to a relative distance of 5, the Nusselt number is again approximately constant and then decreases. It is also significant that the average Nusselt number of line A (maximum value) for the in-line arrangement is always about 20% larger than that of line B (minimum value). In the staggered arrangement, however, no remarkable differences between the Nusselt numbers can be seen for the two lines. Therefore, although the difference between the two arrangements indicates that heat transfer for the staggered arrangement is much more uniform than for the in-line arrangement, the average heat transfer is approximately the same in both arrangements. In Fig. 1.13, the average heat transfer coefficient for a nozzle array is shown as a function of the discharge velocity for selected nozzle diameters. It can be seen that 12 1 Impinging Jet Drying

Fig. 1.13 Average heat transfer coefficient in a nozzle array.

it is possible for a nozzle array to achieve very high heat transfer coefficients in comparison to other forms of flow.

1.3.2 Hole Channels

Nozzle arrays are technically easier to manufacture in the form of hole channels than in the form of individual nozzles; such a hole channel is shown, in principle, in Fig. 1.2b. In the rectangular channel, the holes are punched in the smaller lateral surface and serve as nozzle openings. In this case, the resistance of the outflow must be significantly higher than for the axial flow in the channel to ensure that the same amount of air emerges through all holes. This is usually the situation if the cross- sectional area of the channel is about 2.5-fold greater than the sum of the areas of the perforations. Uneven punching, for example by nonuniform ridges, will result in the jet streams not being perpendicular to the surface, and this in turn will lead to an uneven distribution of heat transfer. In Fig. 1.14, the average heat transfer is shown as a function of the pitch of the perforations, for example for a Reynolds number with relative distance as a parameter. Regarding other Reynolds numbers in principle gives similar curves. Up to a pitch of about 6 the heat transfer will remain substantially constant, but after this point it will decrease. There is, therefore, no pronounced maximum at a 1.3 Nozzle Fields 13

Fig. 1.14 Average heat transfer in a hole channel. pitch of 6, as occurs with a field of individual nozzles, although at a smaller pitch the heat transfer will remain identical. Nonetheless it is recommended that, even with hole channels, a pitch is selected that is no less than 6 in order to minimize the flow rate, and thus the fan performance. The heat transfer will again be highest for relative nozzle distances in the range 2 < h/d < 5, but will then decrease continuously with increasing distance. The Nusselt functions for the stagnation point and the average value of the field can be approximated by (Attalla, 2005):

0:5 0:4 NuSt ¼ 0:47 Re Pr ; ð1:16Þ Nu ¼ 0:10 Re0:67 Pr0:4: ð1:17Þ

Based on comparisons with the corresponding functions for the single-nozzle array, it is evident that the heat transfer in hole channels is less, by about 35%.

1.3.3 Perforated Plates

In perforated plates, spent holes or slots must be installed so that the injected air can flow out again. The size and number of these outlets depends on the manufacturer of the perforated plates, and consequently various cross-flows can occur between the target area and the plate. How this affects the heat transfer is not yet known. Nonetheless, Martin (1977) developed a Nusselt function with which the heat transfer for a majority of perforated plates can be approximated, by providing the following correlation for aligned and hexagonally arranged perforations:

"# : 6 0 05 : : h F 1 2:2F Nu ¼ Re0 67 Pr0 42 1 þ F : ð1:18Þ d 0:6 1 þ 0:2hðÞ=d 6 F

Here, h/d is again the relative nozzle distance, and F is the square-root of the area ratio of the nozzle openings to the total area: 14 1 Impinging Jet Drying

Fig. 1.15 Influence of the pitch and the nozzle distance on the heat transfer for perforated plates.

sffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi A F ¼ nozzles: ð1:19Þ Atotal

For plates with equally distributed pitch t, it follows that rffiffiffi p d F ¼ : ð1:20Þ 4 t The influence of the relative distance and the nozzle pitch is shown in Fig. 1.15, where heat transfer is seen to decrease continuously with the relative distance. The maximum heat transfer occurs in the case of pitches in the range of 3 to 5, depending on the distance. The heat transfer is lower than that of an array of individual nozzles. As the maximum occurs at a smaller pitch than 6 (as for single- nozzle arrays), a higher flow rate must also be applied. Geers et al. (2008) have largely confirmed the correlation of Martin. Heikkil€a and Milosavljevic (2002) concluded that the correlation provides slightly too- high values only when air temperature is above 400 C. Huber and Viskanta (1994) provided a somewhat more simple correlation for pitches t/d 4rightof the maximum (see Tab. 1.2), whereby the Nusselt number was seen to decrease continuously with the pitch.

1.3.4 Nozzles for Cylindrical Bodies

For nonplanar bodies the curvature arises as an additional geometric parameter. In the following discussions, a cylindrical body is considered as the fundamental case. In Fig. 1.16, a cylinder of diameter D is shown, which is impinged from a slot of 1.3 Nozzle Fields 15

Fig. 1.16 Flow to a cylinder from a slot nozzle. width s. With regards to the influence of the geometric parameters and the Reynolds number, divergent results have occasionally been reported in the literature. For example, Gori and Bossi (2003) and Chan et al. (2002) noted that a maximum in the heat and mass transfer occurred at a relative nozzle distance of h/ s ¼ 8, whereas Nada (2006) gave this maximum as between h/d ¼ 4 and 6. McDaniel and Webb (2000) measured a maximum at h/d ¼ 5, but only from nozzles with rounded edges. In contrast, for sharp-edged nozzles the heat transfer was continuously decreased with distance. Olsson et al. (2004) also indicated a decreasing heat transfer with distance, but this was very moderate with an exponent of only 0.077. For all of these authors the influence of the distance was weak. As noted previously for single nozzles and flat surfaces, a distance up to values of h/d ¼ 5 had no effect, whereas heat transfer was slightly decreased at higher distances. The Nusselt functions given below are therefore based on a distance in the range 2 h/s 8. The influence of the ratio D/s has been reported by the various authors with exponents ranging from 0.03 to 0.22; consequently, in order to provide a better comparison their results were approximated with an average exponent of 0.1. The exponents for the Reynolds number specified by the above-mentioned authors have values from 0.4 to 0.82, depending partly on the geometric size. In Fig. 1.17, the measured Nusselt numbers of various authors are compared using D/s ¼ 2andh/s¼ 5 as examples. The length of the dotted lines represents the range of the investigated Reynolds numbers. The bold line with the gradient 0.67 represents an average value; this line can be approximated in 16 1 Impinging Jet Drying

Fig. 1.17 Nusselt number dependent on the Reynolds number for cylinders with slot nozzles for D/s ¼ 2 and h/s ¼ 5.

the range 2 h/s 8with

¼ : 0:67 0:4ðÞ= 0:1: ð : Þ NuD 0 20 ReD Pr D s 1 21 The dimensionless numbers are formed with the cylinder diameter D: a ¼ D ; ¼ w D : ð : Þ NuD l ReD n 1 22 The values of the Nusselt numbers are similar to those resulting from the Nusselt function corresponding to cross-flow of cylinders if the cross-flow length ðÞp D =2 is used as a characteristic dimension in the Nusselt and Reynolds numbers. The Nusselt function can alternatively be formed with the slot width s as the characteristic dimension resulting in

¼ : 0:67 0:4ðÞ= 0:23; ð : Þ Nus 0 15 Res Pr D s 1 23 with a ¼ s ; ¼ ws : ð : Þ Nus l Res n 1 24

1.4 Summary of the Nusselt Functions

The Nusselt functions for the single nozzle are summarized in Tab. 1.1, and for the nozzle fields in Tab. 1.2. The Nusselt and Reynolds numbers are defined with the inner nozzle diameter d regarding round nozzles, and with the width s regarding slot nozzles. The exponent of the Reynolds number in correlations for the average 1.5 Design of Nozzle Field 17

Tab. 1.1 Nusselt functions for single nozzles.

Nozzle shape Nusselt function Scope Reference

Circular Stagnation point region r/d Attalla and Specht ¼ : 0:5 0:4 ; = Nud 0 72 Red Pr 0.5 2 h d 5 (2009) ¼ : 0:5 0:4 = Slot Nus 0 70 Res Pr 2 h s 5 Vader et al. (1991) Circular Averaged values 0 < r/d < 3 Adler (2004) ¼ : 0:67 0:4 Nud 0 12 Red Pr

Tab. 1.2 Nusselt functions for nozzle fields impinging flat surfaces.

Arrays of individual nozzles (in-line or staggered): = ¼ : 0:5 0:4 Stagnation point r d 1: Nud 0 82 Red Pr Attalla and Specht (2009) ¼ : 0:67 0:4 Average: Nud 0 16 Red Pr Pitch t/d ¼ 6, Distance 2 h=d 5 Hole channels: = ¼ : 0:5 0:4 Stagnation point r d 1: Nud 0 47 Red Pr Attalla and Specht (2009) ¼ : 0:67 0:4 Average: Nud 0 10 Red Pr Pitch t/d 6, Distance 2 h=d 4; Width sK=d 2 "# Perforated plates: 0:05 h F 6 1 2:2F Nu ¼ Re0:67 Pr0:42 1 þ F Martin (1977) d d d 0:6 1 þ 0:2hðÞ=d 6 F rffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi rffiffiffiffi A p d F ¼ nozzles, FðÞ¼ round nozzle Atotal 4 t Pitch 1.4 t/d 14, Distance 2 h/d 12 ¼ : 0:67 0:4ðÞ= 0:123ðÞ= 0:725 Nud 0 43 Red Pr h d t d Huber and Viskanta (1994) Pitch 4 t/d 8, Distance 0.25 h/d 6 heat transfer in nozzle fields has various reported values ranging from 0.66 to 0.72. It is mostly a matter of choice which exponent is used for the approximation of data within the spread of the measurement results. In order to enhance comparability of the Nusselt functions, the measurement results of the various references were respectively approximated by the exponent 0.67.

1.5 Design of Nozzle Field

Based on the given Nusselt functions for the required heat transfer coefficient in Eq. 1.9, nozzle arrays can now be designed. It is again assumed that the heat is predominantly transferred for evaporation, while the enthalpy to heat up the dry material is again neglected. First, an array of individual nozzles will be considered for which a distinct maximum in the heat transfer results at a pitch of t ¼ 6d. Thus, 18 1 Impinging Jet Drying

Fig. 1.18 Specific energy consumption, air, and product temperature as a function of the drying rate.

for the heat transfer coefficient from Eq. 1.15 with Pr ¼ 0.7, is obtained:

l : : a ¼ 0:14 w0 67d 0 33: ð1:25Þ n0:67 The material properties are to be formed with the average temperature ðÞT0 þ TS =2. The specific energy consumption thus depends only on the three _ parameters – mv, d, and w – the influence of which must be calculated numerically. In Fig. 1.18, the specific energy consumption, as well as the air and product temperatures, are shown as a function of the drying rate. As an example, a velocity of 50 m s 1, a nozzle diameter of 5 mm, and an ambient temperature of 20 C was used. The higher the drying rate, the more the air must be heated to achieve the enthalpy of evaporation because the velocity, and therefore the flow rate, remain constant. The air does not need to be heated for drying rates less than about 2gm 2 s 1, as an ambient air temperature of 20 Cissufficient. Consequently, the specific energy consumption will be zero. At higher drying rates, the specific energy consumption is initially increased sharply and then passes through a maximum; however, the specific energy consumption decreases steadily thereafter. In Fig. 1.19, specific energy consumption is shown separately as a function of the evaporation rate, with the nozzle diameter as a parameter. The nozzle exit velocity is assumed to be constant at 50 m s 1. It is evident that the specific energy consumption is less with smaller nozzles, with the maximum being displaced towards higher evaporation rates. However, the lowest possible value of nozzle diameter is limited, first by the increase in production costs, and second through the minimum distance of the nozzle to the product. This distance should not be greater than approximately five-times the nozzle diameter; otherwise the heat transfer will be decreased. Rather, the smaller the nozzle the closer the field must be shifted towards the product. In Fig. 1.20, specific energy consumption is shown as a function of the drying rate, with the nozzle discharge velocity as a parameter; as an example, a nozzle 1.5 Design of Nozzle Field 19

Fig. 1.19 Specific energy consumption as a function of the drying rate for different nozzle diameters.

Fig. 1.20 Specific energy consumption as a function of the drying rate for different air velocities.

diameter of 5 mm is used. Here, the lower the discharge velocity, the lower is the specific energy consumption, and the maximum is shifted towards lower drying rates. Lower velocities have the further advantage that only a small pressure drop, and therefore less fan power, will be required. However, the gas (i.e., the air) must be heated to higher temperatures in order to transfer the heat required. In Fig. 1.21, the heating temperature of the nozzle flow is shown as a function of the drying rate, with the nozzle exit velocity as a parameter; the nozzle diameter is again 5 mm. The required temperature increases approximately linearly with the drying rate, and the lower the velocity the steeper the gradient. The lowest value of the velocity is thus limited by the maximum temperature of the gas flow. This maximum temperature is dependent on the type of apparatus used for heating, the material strength, and the thermal sensitivity of the product. Due to the strength of the steel material the gas temperatures are generally limited to 500 C. 20 1 Impinging Jet Drying

Fig. 1.21 Influence of the drying rate and the velocity on the required air temperature.

Fig. 1.22 Influence of the nozzle diameter.

Figures 1.22 and 1.23 show the specific energy consumption, the product temperature and the gas temperature as a function of the nozzle diameter and of the discharge velocity, respectively, as an example for the drying rate 10gm 2 s 1. All three parameters decrease with decreasing nozzle diameter, while the gradients become larger with smaller nozzle diameters. With decreasing velocity the specific energy consumption also decreases and both temperatures are increased. The lower the velocity, the larger is the gradient of variation of the considered parameters. For hole channels, analogous results are valid, and a pitch of t ¼ 6d is again recommended. A larger pitch will result in a decrease in the heat transfer, which will remain constant with lower pitches; however, the number of nozzles and thus the flow rate, will be increased. The specific energy consumption and the required gas temperatures are slightly higher for hole channels than for single-nozzle arrays, because the heat transfer is somewhat lower. 1.5 Design of Nozzle Field 21

Fig. 1.23 Influence of the air velocity.

Fig. 1.24 Comparison of single-nozzle array and perforated plate for different drying rates.

In the case of perforated plates the heat transfer is somewhat less but they are the easiest to manufacture and consequently are commonly used. In Fig. 1.24, the specific energy consumption, the air temperature and the temperature of the material are again shown as a function of the drying rate, with a pitch of t ¼ 6d as an example. The values for the perforated plate are compared here with values for the single-nozzle array. It is clear that, in order to achieve the same drying rate with perforated plates, higher air temperatures – and thus higher specific energy consumptions – are required to compensate for the reduced heat transfer. Figure 1.25 shows the influence of the diameter on specificenergy consumption and both temperatures, using examples of drying rates of 10gm 2 s 1 andof50gm 2 s 1; the pitch is again t ¼ 6d. When the drying rate is 10 g m 2 s 1,thespecific energy consumption of the perforated plate is almost 50% higher than that of the single-nozzle array. However, at a higher drying rate of 50 g m 2 s 1 for diameters larger than 3 mm, the material can no 22 1 Impinging Jet Drying

Fig. 1.25 Influence of the nozzle diameter for two drying rates.

longer be dried; otherwise, the air temperatures must be above 500 C. With regards to perforated plates, the heat transfer according to Fig. 1.15 also depends on the pitch and nozzle distance. Thus, the maximum heat transfer will occur with a pitch in the range of four, which is a smaller value than for single-nozzle arrays. Heat transfer will also depend on the nozzle distance, and increases with decreasing distance. The influence of the pitch is shown for perforated plates in Fig. 1.26, where the pitch is with four smaller than before and with eight greater than before, while the diameter and velocity are held constant. With regards to the pitch of four, lower air temperatures will be needed as the heat transfer will be stronger, although the flow rate will be increased due to the larger number of perforations and, as a result, the specific energy consumption will also increase. In terms of low energy consumption, a large pitch is desirable; however, the maximum possible drying rates are limited duetoincreasingtemperatures. Finally, the influence of the nozzle distance is shown in Fig. 1.27 where, the lower the distance, the lower is the energy consumption. For a given air temperature with lower distances, higher drying rates can be achieved. 1.6 Conclusion 23

Fig. 1.26 Influence of the pitch for perforated plates at different drying rates.

Fig. 1.27 Influence of the nozzle–surface distance for perforated plates.

1.6 Conclusion

The highest convective heat transfer coefficients in drying processes can be achieved using nozzle fields, and drying is strongly intensified in such a field. As a consequence, the heat transfer area, and therefore the size of the apparatus, can be reduced, saving investment costs. The energy required to heat the drying air decreases with smaller nozzle diameters and slower outflow velocities, and consequently the operation costs arereduced.Thelowestpossiblevalueofthenozzlediameteris,however, limited to ensure that the distance between the nozzle and material is not below four diameters. The lower the air velocity, the higher the hot air temperature must be. The temperature of the hot air is limited to 500 Cdue to the strength of steel; however, the hot air temperature must normally be kept much lower in order to protect sensitive products against thermal damage. The pitch of the nozzles has an optimum value of six diameters for 24 1 Impinging Jet Drying

the heat transfer with regards to single-nozzle arrays and hole channels. With regards to perforated plates, the optimum value for heat transfer occurs at pitches in the range of four, while the minimum for the specificenergy consumption is achieved with a nozzle pitch of between eight and ten.

Additional Notation Used in Chapter 1

D diameter of cylindrical drying body m d inner diameter of nozzle m F square-root of nozzle opening area to total area H_ enthalpy flow rate J s 1 h nozzle distance from product m m exponent of Reynolds number s slot width, thickness m t pitch between nozzles m

Subscripts

lo local value s saturation (assumed product condition) st stagnation point/region 0 outlet of heater (entrance of dryer)

References

Adler, W., 2004. Experimentelle Bestimmung channels. Diss., Otto von Guericke des W€armeubergangs€ bei der Prallstromung€ University Magdeburg, Germany. uber€ einen hohen Reynoldszahlenbereich Attalla, M., Specht, E., 2009. Heat transfer mittels Infrarot-Thermografie. Diss., Otto characteristics from in-line arrays of free von Guericke University Magdeburg, impinging jets. Heat Mass Transfer 45(5): Germany. 537–543. Angioletti, M., Di Tommaso, R. M., Nino, Chan, T. L., Leung, C. W., Jambunathan, K., E., Ruocco, G., 2003. Simultaneous Ashforth-Frost, S., Zhou, Y., Liu, J. H., 2002. visualization of flow field and evaluation Heat transfer characteristics of a slot jet of local heat transfer by transitional impinging on a semi-circular convex surface. impinging jets. Int. J. Heat Mass Transfer Int. J. Heat Mass Transfer 45: 46: 1703–1713. 993–1006. Attalla, M., 2005. Experimental investigation of Geers, L. F. G., Tummers, M. J., Bueninck, T. heat transfer characteristics from arrays of J., Hanjalic, K., 2008. Heat transfer free impinging circular jets and hole correlation for hexagonal and in-line arrays References 25

of impinging jets. Int. J. Heat Mass Transfer McDaniel, C. S., Webb, B. W., 2000. Slot jet 51: 5389–5399. impingement heat transfer from circular Gori, F., Bossi, L., 2003. Optimal slot height in cylinders. Int. J. Heat Mass Transfer 42: 1975– the jet cooling of a circular cylinder. 1985. Appl. Therm. Eng. 23: 859–870. Mujumdar, A. S. (ed.), 2007. Handbook Heikkil€a, P., Milosavljevic, N., 2002. of industrial drying, CRC Press, Boca Investigation of impingement heat transfer Raton, USA. coefficient at high temperatures. Drying Nada, S. A., 2006. Slot/slots air jet impinging Technol. 20(1): 211–222. cooling of a cylinder for different jets - Huber, A. M., Viskanta, R., 1994. Effect of cylinder configurations. Heat Mass Transfer jet-jet spacing on convective heat transfer 43(2): 135–148. to confined, impinging arrays of Olsson, E. E. M., Ahrne, L. M., Tr€agardh, A. C., axisymmetric jets. Int. J. Heat Mass 2004. Heat transfer from a slot air jet Transfer 37: 2859–2869. impinging on a circular cylinder. J. Food Eng. Martin, H., 1977. Heat and mass transfer 63: 393–401. between impinging gas jets and solid Vader,D.T.,Incropera,F.P.,Viskanta,R., surface, in Advances in heat transfer, (eds J. R. 1991. Local convective heat transfer from Hartnett, T. F. Irvine), Academic Press, New a heated surface to an impinging, planar York, USA, pp. 1–59. jet of water. Int. J. Heat Mass Transfer 34: 611–623. 27

2 Pulse Combustion Drying Ireneusz Zbicinski, Tadeusz Kudra, and Xiangdong Liu

2.1 Principle of Pulse Combustion

Pulse combustion is an advanced technology of significant potential in thermal processing, since it is claimed to increase productivity, maximize utilization of the input fuel, and reduce pollutant emissions in comparison with conventional (continuous) combustion (Kudra et al., 1994; Zbicinski et al., 2002; Mujumdar and Wu, 2004; Kudra and Mujumdar, 2007). Because of enhanced momentum, heat and mass transfer, pulse combustion drying has been identified by the European Process Intensification Center (EUROPIC) as one of 60 promising process intensification technologies. To make it available to drying professionals, the technology report on pulse combustion drying of EUROPIC (code 4.1.4) has been reproduced in the Drying Technology Journal (Kudra, 2008). The term “pulse combustion” originates from the intermittent (pulse) combus- tion of solid, liquid, or gaseous fuel in burners of special design, consisting basically of the combustion chamber and the diffuser, also termed the tailpipe. Such periodic combustion generates intensive pressure, velocity and, to a certain extent, temperature waves propagated from the combustion chamber through a tailpipe to the process volume (an applicator) such as a dryer, calciner, or incinerator. Due to the oscillatory nature of the momentum transfer, pulse combustion intensifies the rates of heat and mass transfer thus affects greatly such processes as thermal drying. The mechanism behind the operation of a pulse combustor is a complex interaction between an intermittent combustion process and pressure/velocity waves that are propagated from the combustor (Fig. 2.1). The process of pulse combustion is initiated when air required for combustion and fuel in the form of a gas jet or a liquid spray are admitted to the combustion chamber to make-up an explosive mixture, which is ignited by a spark plug and burns instantly in an explosion-like manner. At this moment, the air and fuel inlet ports are closed, either by mechanically operated valves or due to the hydrodynamic action of the rapidly rising pressure. The combustion-generated pressure forces the combustion products to flow out through the tailpipe to the process volume. As the hot flue

Modern Drying Technology Volume 5: Process Intensification, First Edition. Edited by Evangelos Tsotsas and Arun S. Mujumdar. Ó 2014 Wiley-VCH Verlag GmbH & Co. KGaA. Published 2014 by Wiley-VCH Verlag GmbH & Co. KGaA. 28 2 Pulse Combustion Drying

Fig. 2.1 Operation of a pulse combustor with flapper valves.

gases discharge, the resulting outward momentum causes the pressure in the combustion chamber to decrease (Fig. 2.2). When the pressure reaches its minimum, the inlet ports open and admit fresh fuel and air into the combustion chamber. This new charge ignites itself due to contact with remnants of hot flue gases left in the tailpipe from the preceding cycle that reenter the combustion chamber during the minimum pressure period. These combustion cycles repeat themselves at a definite frequency, which depends on the design of the pulse combustor and characteristics of the tailpipe. In addition, the startup fan is normally shut off once self-aspiration is established. Typically, pulse combustors operate at frequencies from 20 to 250 Hz. Pressure oscillation in the combustion chamber of 10 kPa produces a velocity oscillation in the tailpipe of about 100 m s 1, so that the instantaneous velocity of a gas jet at the tailpipe exit varies from 0 to 100 m s 1 (Keller et al., 1992). The amplitude of the pressure rise may vary from 10% (domestic heating applications) to 100% as for 2.1 Principle of Pulse Combustion 29

Fig. 2.2 (a) Theoretical and (b) experimental C: combustion complete, pressure decreases pressure trace in a pulse combustor. A: air and as flue gases are vented; D: momentum of fuel enter the combustion chamber; B: fresh exhausting gases creates negative pressure in charge ignited, pressure rises as combustion the combustion chamber. gases heat up, air and fuel inflow are stopped; heavy-duty pulse combustors for industrial use (Kentfield, 1993). The output power for commercially available pulse combustors ranges from 70 to 1000 kW. Numerous studies on a variety of pulse combustor designs have demonstrated that pulse combustion offers numerous benefits over continuous combustion, of which the most important are (Stewart et al., 1991; Keller et al., 1992; Ozer, 1993; Kudra, 2008):

increased heat and mass transfer rates (by a factor of 2 to 5); increased combustion intensity (by a factor of up to 10); higher combustion efficiency with low excess air values; reduced pollutant emissions (especially NOx, CO, and soot) by a factor of up to 3; 30 2 Pulse Combustion Drying

Fig. 2.3 Typical emissions from pulse combustor (courtesy of Novadyne Ltd, Hastings, ON, Canada).

improved thermal efficiency by up to 40%; and reduced space requirements for the combustion equipment.

An added benefit of pulse combustion is its contribution to environment protection. The rapid combustion, which allows extremely short times for the formation of nitrogen oxides, and lower peak temperatures as compared to

continuous combustion, result in very low NOx emissions from the pulse combustors (Fig. 2.3). This is especially advantageous when drying foods and biomaterials, provided that the pulse combustor is not fed with sulfur-containing fuels such as propane or natural gas (Kudra, 1998; Kudra, 1999). Recent studies on pulse combustion have indicated that the design of the combustion chamber and the geometry of a combustor–dryer system may affect the level of noxious emissions. Noticeably lower concentrations of CO, NO and

NOx were obtained for those combinations of the volume of the combustor chamber and length of the tailpipe that provide smoother and sinusoidal pressure fluctuations (Zbicinski et al., 1999). The influence of the tailpipe length and fuel fl ow rate on NOx emission is presented in Fig. 2.4. A higher residence time in longer tailpipes results in increased NOx emission, and the same effect is observed for the fuel flow rate. During measurements performed at optimal geometry and

operating conditions, the NOx level ranged between 5 and 98 ppm and the CO emission between 3 and 40 ppm. When the air-to-fuel ratio was lower than the stoichiometric ratio, leading to incomplete combustion, the CO emission was over 20 000 ppm (Smucerowicz, 2000). The results of experiments showed that the pulse combustor could function over a wide range of air-to-fuel ratios, but that the optimal range from the low-emissions viewpoint was narrow. With regards to other gaseous emissions in drying applications, the high temperatures of combustion gases when in contact with the wet material fed into the tailpipe may result in the destruction of odor, toxic contaminants, and other 2.1 Principle of Pulse Combustion 31

Fig. 2.4 Effect of tailpipe length and fuel flow rate on NOx emission. volatile organic compounds (VOCs), either by thermal dissociation or by free- radical oxidation reactions (Rafson, 1998; Guy et al., 1997). From another viewpoint, concerns of odors are lessened in pulse combustion dryers where the wet material is injected directly to the tailpipe; this occurs as the result of an extremely short residence time (fractions of a second), as well as the effect of evaporative cooling which keeps the product temperatures sufficiently low so as to prevent the release of volatile substances. The major disadvantage of pulse combustion systems is the extreme noise generated under normal operation, which may be a deterrent factor in industrial applications. The sound pressure level (SPL) depends on the operating pressure and frequency; the higher the operating pressure, the higher the SPL. From a technical viewpoint, large pressure amplitudes are desirable in processes such as drying because of increased rates of heat and mass transfer, and thus a reduced size of the burner with the same heat output. From a practical viewpoint, however, noise emission may be a much more important design criterion than minimizing the burner size. Although the SPL at the tailpipe outlet attains 110 dB (Zbicinski, 2002), this noise is greatly attenuated in a drying chamber, with the major source of emitted noise occurring at the inlet to the combustor. As indicated by Duncan (2012), building the combustor from thick-walled Inconel and encasing the entire combustor in an outer shell will reduce the noise level to less than 85 dB. Mounting a sound suppressor built as a tube-in-tube segment, where the in-between tubes volume is filled with a porous material, will reduce the sound level from 160 dB to 140 dB and 90 dB, respectively, when the suppressor is attached at the tailpipe outlet and the combustor inlet. 32 2 Pulse Combustion Drying

Fig. 2.5 Design principle of pulse combustor with rotary valve (after Lockwood, 1987).

Another solution to the sound problem is a rotary-valve combustor enveloped by a steel shell, as shown in Fig. 2.5. Because, for spray-drying applications, such a combustor is running out of its resonant frequency, the sound level will be less than 85 dB at a distance of 1 m from the dryer (Rehkopf, 2012).

2.2 Pulse Combustors: Design and Operation

Pulse combustors may be categorized into three distinct classes according to the specific acoustic system on which their operation depends: (i) the Quarter-wave (or Schmidt) combustor; (ii) the Helmholtz combustor; and (iii) the Rijke-type combustor. In contrast to the Rijke combustor, which operates with solid fuels, both the Schmidt and Helmholtz combustors accept liquid and gaseous fuels. The Helmholtz combustor is preferred for drying applications because the larger volume of the combustion chamber and the smaller (but longer) tailpipe allows for multivalve assembly. Detailed information on these types of combustor is available in articles by Zinn (1985) and Kudra and Mujumdar (2009). Some combustors also exploit the resonance phenomenon; these are referred to as frequency-tunable pulse combustors.

2.2.1 Pulse Combustors with Mechanical Valves

Based on the manner in which fuel and air charge the combustion chamber, pulse combustors can be divided into two general categories: (i) those with mechanical valves; and (ii) those with aerodynamic valves (also called “valveless” combustors). 2.2 Pulse Combustors: Design and Operation 33

Fig. 2.6 Mechanical valves. (a) Flapper-type; (b) Reed-type with petal-shaped lamella; (c) Reed- type with rectangular lamella.

Mechanical valves can be further divided into three subtypes of flapper valves, reed valves, and rotary valves. Since mechanical valves provide a physical barrier to the backflow of combustion products through the combustor inlet during the positive-pressure phase of the pulse combustion cycle, the unidirectional flow is the fundamental feature of valved pulse combustors. There are, however, certain problems associated with the design of mechanical valves, such as minimizing valve inertia, protection from corrosion, and resistance to material fatigue due to thermal stress. These specific problems are of major importance in heavy-duty pulse combustors operated at large pressure amplitudes (Kentfield, 1993). The basic consideration in the design of the flapper (membrane) valves, as shown schematically in Fig. 2.6a, is the ability to rapidly open and close the openings in the base plate according to the combustor operating frequency. Moreover, the valves must withstand the maximum pressure in the combustion chamber, yet open with only a small pressure difference. The reed valves normally used in heavy-duty pulse combustors are made from thin-sheet spring-steel (Fig. 2.6b and c), and the spring action of reed valves is such that, when normally shut, the valves are sprung lightly. In order to ensure a vigorous mixing of the fuel and air, the fully open flow area of the inlet reed valves must be considerably smaller than the cross-sectional area of the combustion zone. The major problem often encountered with reed-type mechanical valves is fatigue- based failure. A schematic diagram of a pulse combustor with flapper valves is shown in Figure 2.7. In contrast to flapper and reed valves, the operation of rotary valves is independent of the combustor design, so they can accommodate a broad range of combustor frequencies and firing rates. Other important features of rotary valves include durability, flexibility of use, and resistance to oil and dirt accumulation. Generally, the rotary valve is built on two plates: a motor-driven “butterfly- shaped” rotating plate; and a stationary plate with two (or more) slots displaced by 34 2 Pulse Combustion Drying

Fig. 2.7 Pulse combustor with flapper valves (after Kitchen, 1987).

180 (or less, depending on the number of slots in the stationary plate). The air for combustion passes through the valve, perpendicular to the rotating plate, and enters the combustion chamber. The height and width of the slots on the stationary plate determine the open area for airflow through the rotary valve; alternatively, the plates may have circular or other shaped openings. The second design of rotary valve is based on three coaxially disposed sleeves with a plurality of slots in their lateral surface (c.f., Fig. 2.5). When rotated by a motor with controllable speed, the slots periodically admit the intake air at the rate required for optimum combustion, whereas the fuel is typically injected via the nozzle. In such a design, as the rotation of the valve determines the running frequency of the pulse combustor, an adequate “feedback system” is required to synchronize the rotation of the valve with the resonant frequency of the burner. The speed of the valve, which influences the resonant frequency, is determined through the sensing and processing of pressure pulses in the combustion chamber.

2.2.2 Pulse Combustors with Aerodynamic Valves

Aerodynamic valves employ the properties of the fluid entering a specially designed inlet to the combustor to create an artificial barrier to the backflow of combustion products out of the combustor through its inlet section. The main advantage of aerodynamic valves is a lack of mechanical parts that are prone to failure. The design concept of aerodynamic valves exploits the principle of a fluid diode which are, of course, inferior in performance when compared to mechanical valves as the backflow of combustion products cannot be fully eliminated. One mechanism that is known to limit the amount of backflow is to include a nonuniform cross- sectional area. A tapered inlet that diverges gradually towards the combustion chamber initially accelerates the inlet stream of air and then diffuses it before 2.2 Pulse Combustors: Design and Operation 35

Tab. 2.1 Operational characteristics of pulse combustor fed with various fuels (Wu, 2007).

Parameter Propane Ethanol Fuel oil

Frequency, Hz 116 120 112 Pressure amplitude, Pa 5082/þ6035 6351/þ6800 5283/þ6404 Average temperature in chamber, K 2094 2106 2070 Average gas velocity, m s 1 16.55 15.69 17.67 Fuel flow rate, g s 1 0.1929 0.3077 0.2162 Combustion heat, MJ kg 1 46.36 28.0 40.53 Power output, kW 8.95 8.64 8.76 Average trust, N 0.4244 0.4018 0.4595 Specific impulse, Ns kg 1 2200 1310 2130 Trust/power output, N kW 1 0.0474 0.0465 0.0524

Fig. 2.8 Configuration of University of Calgary pressure-gain valveless pulse combustor (courtesy of J. A. C. Kentfield). entering the combustion chamber. Such a tapered inlet acts effectively as a nozzle restricting the backflow of combustion products. Figure 2.8 presents, as an example, the schematics of a reverse-flow, pressure-gain, valveless pulse combustor designed at the University of Calgary. Details for the designs of valves and various pulse combustors can be found elsewhere (Kentfield, 1993; Speirs, 1989; Olorunmaie and Kentfield, 1989). The performance characteristics of a flapper valve pulse combustor fed with various fuels are listed in Tab. 2.1 (Wu, 2007).

2.2.3 Frequency-Tunable Pulsed Combustors

Pulsed combustors usually operate at a frequency which is much lower than the inherent frequency, often controlled by an ignition, fuel injection, or a valve open/ 36 2 Pulse Combustion Drying

close sequence. Thus, valveless or flapper valve combustors fall into a category of pulse combustors while mechanically driven valves (e.g., rotary valve) used to control either air or fuel inflow, flue gas discharge (or both) are categorized as frequency-tunable pulse combustors. One possible design of a frequency-tunable pulse combustor takes advantage of the natural non-longitudinal acoustic modes of the process chamber, such as a dryer or incinerator (Zinn and Daniel, 1988). In order to obtain maximum benefit from the pressure/velocity oscillations in the process chamber, the pulse combustor can be tuned to one (or more) of these acoustic modes by modulating the flow of fuel to the combustion chamber. Such modulation may be accom- plished by exciting the acoustic resonance within the fuel line, or via periodic interruptions of the fuel flow, using a rotary valve. The frequency-tunable pulse combustor combines a combustion zone and exhaust zone. The reaction between fuel and air, which takes place in the combustion zone, induces an acoustic wave in the combustor that is then used to control the flow of fuel and air into the combustion chamber (Lockwood, 1987). The acoustic characteristics of the combustor, which may deliberately be altered to provide a selectively variable (or frequency-tunable) pulsating combustion, are briefly described by Kudra and Mujumdar (2009).

2.3 Aerodynamics, Heat and Mass Transfer

Since transport phenomena in the pulse combustion chamber are beyond the scope of this book, the following sections will be restricted to momentum, and to heat and mass transfer between the combustion gases discharged through thetailpipeandtheparticlesofadrying material or droplets of atomized liquid. Even though the microscopic effects in pulse combustion drying have been well identified and exploited in industrial dryers owing to the professional experience of dryer manufacturers and field testing (e.g., Rehkopf, 2012; Duncan, 2012), an insufficient understanding of the process mechanism and a scarce mathematical description of the transfer phenomena (Wu and Mujumdar, 2006a; Xiao et al., 2008; Liewkongsataporn et al., 2008) has hampered the science-based dryer design and scale-up. Aside from exploratory research by Kudra et al. (1994) to compare drying rate of solid wood in pulsed and continuous combustion, as well as recent studies on the pulse combustion drying of a paper sheet (Liewkongsataporn et al., 2006; Liewkongsataporn et al., 2008), the majority of publications have described the pulse combustion drying of dispersed materials (Kudra and Mujumdar, 2009). Thus, the following sections will provide an overview of the investigations conducted on transfer phenomena in droplets and granules exposed to the oscillating flow of pulse combustion gases. 2.3 Aerodynamics, Heat and Mass Transfer 37

2.3.1 Atomization

Following on from the aforementioned characteristics, an advantage of pulse combustion is that the strongly oscillating hot-gas jet from the combustion chamber, and also from the tailpipe, can promote the dispersion of liquids, pastes and clumps of particulate materials (Wu et al., 2012). Hence, in the case of liquids, pulse combustion spray-dryers can refrain from using conventional disk or nozzle atomizers, which translates into a simple design and lower capital and operating costs, as well as an enhanced product quality due to the creation of finer droplets with a narrow size distribution. The process of liquid atomization by the oscillating gas jet emerging from a Helmholtz-type pulse combustor was studied experimentally for water and aqueous maltose solutions with different viscosities (Xiao et al., 2008). The frequency of the oscillating flow of combustion gases ranged from 61 to 100 Hz, while the pressure spanned from 93.7 to 114.8 kPa with an average amplitude of about 21 kPa. Because temporal variations of temperature and velocity were difficult to measure for such high pulse frequency fluctuations, the cycle-averaged values of air flow temperature and velocity were used in the data interpretation. The air temperature in the tailpipe was maintained at 423 10 K using a water-cooled system. The oscillating air flow transient velocity was between 100 m s 1, and the mean velocity about 26.8 m s 1. The droplet size was presented in terms of Sauter mean diameter (SMD), which is the diameter of a droplet having the same volume-to- surface area ratio as the ratio of the total volume of all droplets to the total surface area of these droplets. Figure 2.9 presents the effect of liquid feed rate, liquid viscosity and oscillating frequency on the SMD of the dispersed liquid. Effects of all three parameters were evident with, most notably, larger droplets being obtained for a higher liquid feed rate. When the liquid feed rate was below 28 l h 1, however, the atomized droplets were so small that they vaporized immediately and the optical analyzer had insufficient time to capture their image, and consequentlyallsubsequenttestswereperformedataliquidfeedrateof 35 l h 1. Smaller droplets were obtained at a higher frequency of gas flow oscillation, with the best atomizing effect being achieved at a liquid feed rate of 35 l h 1 and the gas flow oscillating at a frequency of 100 Hz. The experimental results presented in Fig. 2.9 show that the liquid viscosity will have a major impact on the characteristics of atomization. An inversion point was observed at a viscosity of 0.007 Pas whilst, at a relatively low viscosity (<0.007 Pas), the droplet SMD decreased with increasing liquid viscosity. When the viscosity was above 0.007 PasthedropletSMDwasseentoincreasewiththeviscosity.It should be noted that higher surface tension of the liquid drop may bafflethe atomizing process and result in a larger droplet SMD. Therefore, a better liquid atomization will occur at a moderate liquid viscosity level if all other parameters are held constant. 38 2 Pulse Combustion Drying

Fig. 2.9 Droplet size variation with liquid feed rate, viscosity, and frequency.

2.3.2 Heat and Mass Transfer

The main difference between pulse combustion drying (PCD) and conventional steady-flow drying is the environment for heat and mass transfer between the discrete and continuous phases. In the PCD process, an unsteady, turbulent flow (turbulence level up to 4000%, as shown by Smucerowicz (2000)), gas velocity oscillations (50–200 Hz), and a high temperature difference between the gas and the drying material may have a significant effect on the drying rate and heat and mass transfer coefficients. As an enhancement of the drying rate might be caused by an increase in the heat transfer coefficient, the heat transfer area or/and the driving force of the process, the following factors should be taken into account:

The high driving force of the process: This is the most important factor accelerating the drying rate. The difference between gas and wet material temperatures (which may reach about 700–800 K at the inlet) controls the drying time and results in a rapid and efficient drying process. Stripping of the boundary layer: The heat transfer coefficient can be significantly affected due to strong velocity oscillations and stripping of the boundary layer. An analysis performed by Ozer (1993) and presented in Fig. 2.10 shows a periodical shift of the position of the stagnation point during oscillating flow around a cylinder, which must influence the local and global Nusselt numbers. Secondary atomization: This increases the heat transfer area due to the presence of more fines in the droplet size distribution. Murray and 2.3 Aerodynamics, Heat and Mass Transfer 39

Fig. 2.10 Temporal effect of velocity oscillations on boundary layer (Ozer, 1993).

Heister (1999) confirmed droplet break-up in the oscillating flow and found the break-up times to be roughly inversely proportional to the Weber number. An analysis of the literature confirmedthatthefreejet can atomize slurries with viscosities up to 2 Pas, and solutions with viscosities up to 0.3 Pas (Ozer, 1993).

In the past, numerous groups have attempted to determine the increase in the evaporation rate of PCD as compared to steady convective drying, with most findings suggesting an approximately twofold rate for PCD. However, reports have also been made of significantly different estimations of the effect of velocity oscillations on heat and mass transfer in PCD. Nomura et al. (1989) reported a fivefold higher drying rate in the flow field generated by a pulse combustor than in a steady flow for the drying of bricks, while Kezerle (1993) achieved only a 5% enhancement of heat and mass transfer rate for droplet evaporation. Similar findings were described by Buchkowski (1999) for a pulse combustion flash-drying process (10%). Saito et al. (1996) reported an increase in the evaporation rate coefficient by a factor of up to three compared to a non-oscillating environment for evaporation during the direct combustion of fuel droplets in a Rijke tube. Habib et al. (2002) showed that flow oscillations can enhance the heat transfer coefficient under laminar flow conditions, but this effect proved to be negligible at higher flow rates – that is, in the turbulent regime. Both Akulich and Kuts (2003) and Akulich et al. (2010) showed, on a theoretical basis, how the relative gas/particle velocity and particle size would affect the gas-to-particle heat transfer rate in PCD. For higher gas flowfrequencies,theparticlesareunabletofollow the oscillations of the gaseous carrier, and so the heat transfer coefficient increases with gas pulsation frequency due to a higher velocity difference between the gas and particles. With regards to particle size, smaller particles tend to more closely follow the gas flow; as a consequence their relative velocity decreases and this leads to lower heat transfer coefficientsforsmallerparticles compared to larger particles. Liewkongsataporn et al. (2006) analyzed the drying of paper by the impingement of a pulsating jet discharged from a valveless pulse combustor for different nozzle- to-plate distances and tailpipe diameters. The results of these experiments indicated a twofold enhancement in relation to impingement drying by means of a steady, nonpulsating jet. Similar experiments on the PCD of paper by an impinging jet were performed by Wu and Mujumdar (2006a) and Wu (2007). These authors found that the local heat transfer coefficient varied from 126 W m 2 K 1 at the 40 2 Pulse Combustion Drying

stagnation point to 3 W m 2 K 1 at a radial distance equal to five times the tailpipe diameter. Wu (2007) explained that the high values of heat transfer coefficient at the stagnation point were due to the impinging jet flow in this region, whereas the gradual decrease in gas jet temperature due to a progressively increasing entrainment of ambient air reduced the intensity of heat transfer in a radial direction outward the stagnation point. Liu et al. (2001) placed a 10 mm brass sphere in the flow field inside the tailpipe of a Helmholtz-type pulse combustor with an external flapper valve to determine the convective heat transfer coefficient under various oscillation frequencies. Figure 2.11 presents the results of measurements of sphere temperature and heat transfer coefficient in a gas stream oscillating at 75 Hz. Based on the results of these experiments, Liu et al. (2001) determined the following correlation between the Nusselt number (Nu) and the frequency of oscillations (f) over the temperature range from 300 to 900 C:

: : : Nu ¼ð2:0 þ 0:6Re0 5 Pr0 33Þð1:0 þ f Þ0 15; 3:5

Fig. 2.11 Heat transfer coefficient in a gas stream oscillating at 75 Hz (Liu et al., 2001). 2.3 Aerodynamics, Heat and Mass Transfer 41

: : : Sh ¼ð2:0 þ 0:6Re0 5 Sc0 33Þð1:0 þ f Þ0 29; 3:5

Fig. 2.12 Flux of liquid phase as a function of a distance from the atomizer. 42 2 Pulse Combustion Drying

Tab. 2.2 Percentage of evaporated water.

Distance from the nozzle, cm Atomization ratio (air/water), kg h 1/kg h 1

8.4/5 8.4/7 12/7 8.4/10 8.4/15 18/15

7 89.3 46.9 62.3 91.1 97.1 35.5 17 97.8 79.1 97.1 97.6 98.5 — 37 99.8 97.2 99.1 99.8 99.9 35.5 60 99.9 99.7 99.9 100 100 77.4

Smucerowicz (2000) performed PCD experiments with NaCl solutions for different operating parameters of the valved pulse combustor. Initially, a filter paper technique was used to trap material in the drying chamber and to determine its moisture content. The results for 5% and 10% aqueous solutions of NaCl showed that, for a higher salt content, the drying process was completed more rapidly due to the atomization of these solutions into much smaller droplets. To summarize, in the PCD process convective heat and vapor transfer between the material surface and the drying gas is strongly affected by the flow oscillations of the gas stream. Additionally, sound waves produced by the combustor can penetrate the solid materials, affecting the internal moisture transport rate. The local pressure difference induced by the acoustic field may cause an enhanced removal of water vapor from capillaries, and also promote the movement of liquid water towards the material surface. Scarboroug et al. (2006) observed a significant increase in the transfer coefficients when the ratio of the acoustic velocity amplitude to the mean velocity was greater than 1.8, which means that the acoustic oscillations of an appropriate amplitude could increase the drying rate of the pulse combustion process.

2.4 Modeling of Pulse Combustion Drying

Pulse combustion is a complex phenomenon that has proved difficult for both experimental and numerical analyses. Wu and Mujumdar (2006b), who first analyzed the needs for basic and applied research in PCD, stressed the lack of any mathematical models that could be validated by the measurement of flow and temperature fields, as well as noise levels, for an effective scale-up of the process. Recently, robust developments in the capabilities of computers have led to the modeling of transient turbulent flows becoming much less challenging. The experimental analysis of PCD, which requires sophisticated measurements (e.g., LDA, PDA, pressure, temperature, noise, etc.) is difficult, hostile (e.g., high noise level, around 110–130 dB) and expensive to carry out in comparison with numerical simulations (Zbicinski, 2002). Computational fluid dynamics (CFD) models for a steady or transient flow, for example, as generated by the pulse combustor, differ 2.4 Modeling of Pulse Combustion Drying 43 only in their definition of velocity at the inlet to the drying chamber. Whereas, for a steady flow the inlet air velocity must be constant, a user-defined function must be employed for a pulsating flow in order to calculate the velocity oscillations at the entrance to the dryer. Nonetheless, a few attempts have been made to mathematically model the pulse combustion process (Ponizy and Wojcicki, 1985; Barr et al., 1988; Celik et al., 1993; Morel, 1993; Lindholm, 1995; Kuts et al., 2002 (analytical model); Wu and Liu, 2002; Wu and Mujumdar, 2006a; Wu, 2007), in addition to a small number of validated models (Zbicinski et al., 1999; Smucerowicz, 2000). The only validated CFD model of a pulse combustion spray-drying process has been presented by Zbicinski et al. (1999) and Smucerowicz (2000). The two- dimensional axi-symmetric CFD calculations of water evaporation and drying of salt solutions in the oscillating flow of a continuous phase were carried out for a valved pulse combustor. Extensive experiments (PDA, LDA, temperature, noise, pressure measurements) to determine the inlet conditions and also to validate the model were performed for a horizontal pulse combustion dryer (Smucerowicz, 2000). Sinusoidal axial velocity oscillations have been defined as follows:  uz ¼ u þ uoscil sinðftÞ; ð2:3Þ  where uz is the velocity of a pulsating gas jet generated by the pulse combustor, uis the mean velocity, uoscil describes the amplitude of velocity oscillations, f is the frequency of oscillations, and t represents time. All velocities and the frequency of the pulse combustor were determined experimentally (Smucerowicz, 2000):  ¼ : 1 ¼ 1 ¼ 1 u 5 76 m s ,uoscil 12 m s , and f 754 rad s (121 Hz). In the pulse combustion spray dryer (Wu and Mujumdar, 2006a) the values were as follows:  ¼ 1 ¼ 1 ¼ 1 u 23 m s ,uoscil 57 m s , and f 520 rad s (83 Hz). CFD simulations carried out for steady-state and oscillating flows generated by the pulse combustor allow the comparison of flow fields in the drying chamber. The same boundary and initial conditions (except for the velocity oscillations) were assumed for the steady and transient cases. Figure 2.13 shows a typical streamlines pattern characteristic of a steady spray-drying process (e.g., Langrish and Zbicinski, 1994; Oakley et al., 1988), with two recirculation zones occurring near the inlet and outlet of the drying chamber. Figure 2.14a–d displays the streamlines pattern in the drying chamber for an oscillating gas flow during one cycle after 19.5 s (at 19.5, 19.502, 19.504, and 19.507 s) when a pseudo-stable state was achieved. The main flow is shifted from the dryer axis to the walls, while one recirculation zone is formed near the dryer inlet and another is formed temporarily in the conical section of the dryer as a result of reverse flow. Recirculation zones are responsible for negative values of axial velocity near the wall. The calculated streamlines pattern resembled the flow pattern in a classical spray-drying process. Similar predictions were reported by Wu and Mujumdar (2006a). Figures 2.15 and 2.16 show families of curves representing the velocity and temperature oscillations. Changes in velocity at the dryer inlet induce oscillations in velocity in the entire dryer volume; these oscillations spread out along the dryer length but decrease continuously with increasing distance from the dryer inlet. No 44 2 Pulse Combustion Drying

Fig. 2.13 Streamlines pattern in a spray-dryer (steady flow).

Fig. 2.14 Streamlines pattern in a pulse combustion spray-dryer (pulsating flow) after (a) 19.5 s, (b) 19.502 s, (c) 19.504 s, and (d) 19.507 s.

inlet temperature oscillations have been defined in the CFD model, although due to oscillations in the flow the air temperature is also affected. Following the velocity oscillations, temperature oscillations decrease with dryer length and fade away at 20 cm from the dryer inlet, as illustrated in Fig. 2.16. Similar observations were presented by Wu (2007). Temperature profiles for the evaporation of salt solution (atomization ratio 18/15 kg h 1/kg h 1) are presented in Fig. 2.17. A rapid depression of drying agent temperature near the atomizer, where an intensive evaporation of water takes place (typical also for classical spray drying), can also be observed. In pulse combustion spray drying the evaporation of water to the constrained volume of the spray envelope in the vicinity of the atomizer reduces the temperature of the drying agent in the dryer axis. However, this effect can be observed only near the atomizer where the radius of the spray envelope is small. Further extension of the spray envelope 2.4 Modeling of Pulse Combustion Drying 45

Fig. 2.15 Axial velocity oscillations in the dryer axis.

Fig. 2.16 Temperature oscillations in the dryer axis. allows more fresh air to come into contact with the dispersed phase and, despite evaporation, this results in increasing temperature and flattening gas temperature profiles across the radius of the spray envelope (Fig. 2.17). Due to velocity and pressure oscillations the temperature profiles are equalized faster than in the classical spray drying process. Wu and Mujumdar (2006a) performed an extensive theoretical analysis of the effects of gas temperature, pulse frequency, amplitude and gas mass flow rate on the transient flow patterns, droplet trajectories and overall pulse combustion 46 2 Pulse Combustion Drying

Fig. 2.17 Air temperature profiles for the evaporation of salt solution (atomization ratio 18/15 kg h1/kg h1) in the vicinity of the atomizer.

spray drying performance. These authors confirmed the occurrence of a recirculation zone near the inlet to the drying chamber (c.f., Zbicinski et al., 1999) that oscillated with the pulse combustor frequency (83 Hz) and which promoted intensive mixing and an augmented heat and mass transfer between the disperse and continuous phases. An important result from this model was that droplet drying conditions varied at different injection times over a pulse period due to changing flow patterns in the drying chamber. Figure 2.18 displays the changes in particle moisture content for droplets injected into the chamber at different points of time during the period of 0.254 to 0.259 s. Particles with a moisture content close to the outlet or to the inlet moisture content(2kgkg 1 and 9 kg kg 1, respectively) can be found in the chamber, this being the result of droplets conveyed by a pulsating gas jet into the recirculation zone where they can be dried completely, or to the chamber walls where less moisture is removed. Another important finding from the analysis was a limit in pulse amplitude beyond which no improvement of droplet drying rate was observed. Wu and Liu (2002) analyzed the mechanism of the pulse combustion spray drying of salt solution in an oscillating flow fieldproducedbyaHelmholtz- type pulse combustor with an external flapper valve. The solution was atomized directly by the pulsating flow in the combustor tailpipe, which reduced the average droplet diameter by about 50% in relation to conventional nozzle atomization. An optical analyzer was applied to measure the droplet size distribution in the spray in order to determine the initial conditions for the CFD model of the process. The results of CFD simulations enabled the average residence time of droplets in the drying chamber (ca. 0.1 s) and the 2.4 Modeling of Pulse Combustion Drying 47

Fig. 2.18 Changes of moisture content of droplets injected into the chamber at different points of time between 0.254 and 0.259 s. evaporation rate (which was about sixfold higher than that in traditional spray drying) to be determined. The results of the CFD analysis revealed a large- scale vortex in the upper part of the drying chamber (as in Zbicinski et al., 1999). The model also predicted the effect of high instability of the flow field in the chamber on particle trajectories during one pulsation period. Wu (2007) elaborated a CFD model for parametric studies of a pulse combustor with mechanical valves to explain the pulsating flow characteristics. On the basis of numerical results, a small-sized pulse combustor was designed and tested that provided a good agreement with experimental observations. By using a CFD model, Ahrens et al. (2011) determined the nozzle spacing in the impingement zone during pulse combustion and the steady impingement drying of tissue (“Yankee dryers”). These authors also proved that the jet thermal energy effectiveness may be threefold higher for pulsating impingement jets than for steady-flow jets at the same temperature and mean velocity. The CFD model not only provided valuable information on the mechanism of the PCD process but also enabled the design of a nozzle-to-nozzle configuration to increase the productivity of Yankee dryers. To summarize, the most valuable information obtained to date on the mechanism of pulse combustion spray drying processes is derived from CFD models. An analysis of the results of such models has proved that changes in the basic process parameters of pulse combustion spray drying are similar to these in classical spray drying, with recirculation zones being formed in the inlet part and conical section of the dryer. The streamlines pattern in pulse flow resembles the flow pattern in a classical spray drying process. However, oscillations in the large recirculation zone between the dryer axis and the wall are observed during each 48 2 Pulse Combustion Drying

cycle of combustor operation. As the experimental analysis of the PCD process is known to be difficult, hostile, and expensive, CFD models have become increas- ingly valuable tools not only for delivering information on the mechanism of PCD but also assisting in the design of pulse combustion systems, including the shape of the drying chamber (e.g., Wu, 2007), the increased efficiency of a drying process (e.g., Ahrens et al., 2011; Widiyastuti et al., 2007), and reductions in noise emissions (e.g., Zbicinski et al., 1999). However, a confident and effective scale-up presupposes the development and use of validated CFD models of PCD processes (Wu and Mujumdar, 2006b; Kudra, 2008).

2.5 Pulse Combustion in Drying

Today, PCD can be performed either within the pulsating gas stream in the tailpipe or, more frequently, in a separate cavity (also called an applicator or a processing volume) that functions as a sound resonator. Thus, the typical pulse combustion dryer consists of a combustor and a drying chamber where the hydrodynamic action of pressure and velocity waves enhances the drying rates (Buchkowski and Kitchen, 1993; Buchkowski and Kitchen, 1995; Kudra et al., 1994; Kudra et al., 2003; Mujumdar and Wu, 2004). Because the primary function of a pulse combustor in this configuration is to excite large-amplitude velocity/pressure oscillations, the pulse combustor may supply only a fraction (10–30%) of the heat required for evaporation, which opens retrofit applica- tionstoexistingrotary,belt,spray,andother dryers. Figure 2.19 presents such aretrofit of PC to a rotary dryer (Richards, 1994).

Fig. 2.19 Rotary kiln with frequency-tunable pulse combustor (Richards, 1994). 2.5 Pulse Combustion in Drying 49

Fig. 2.20 Typical configuration of a pulse combustion spray-drying system.

The most advantageous configuration of pulse combustion drying is to disperse a liquid material within a pulsating jet of flue gases (Fig. 2.20) which is then dried in a spray drying chamber (Stewart et al., 1991; Keller et al., 1992; Ozer, 1993; Ozer et al., 1993; Zbicinski et al., 1998; Rehkopf, 2012). The power output of one of the commercially available units is 900 kW, which allows the evaporation of up to 900 kg of water per hour with 1800 kg of drying air per hour (Rehkopf, 2012). The dryer can handle not only viscous liquids but also slurries; a pulsating jet of flue gases can atomize a slurry containing suspended solids up to 2 mm in diameter and up to 10 mm in length; moreover, there is no need to apply a nozzle to atomize the slurry. The particles of the product are smaller than in a spray dryer using conventional nozzles, and vary from 30 to 60 mm in size. A narrower particle size distribution is of interest to industries dealing with nanomaterials, metals, minerals and ceramics, whereas gentle atomization and a short residence time allows the drying of biomaterials such as cyclodextrin (Rehkopf, 2012). The basic characteristics of classical and pulse combustion spray dryers are compared in Tab. 2.3. 50 2 Pulse Combustion Drying

Tab. 2.3 Basic differences between classical and pulse combustion spray dryers (courtesy of Pulse Combustion Systems, Payson, AZ, USA; www.pulsedry.com).

Dryer characteristics PCS Pulse Dryer Spray dryer (nozzle Spray dryer (disk distributor) distributor)

Atomization Gas dynamic, no High-pressure orifice, Moderate shear shear high shear Feed pump Low pressure High pressure Low pressure (7 kPa) Feed delivery to dryer Open pipe Precision orifice Various wheel types Evaporative heat At point of Remote, ducted to Remote, ducted to atomization drying chamber drying chamber Mixing of hot air and Rapid and Slow Slow droplets complete Heat transfer rate High Relatively low Relatively low Parts that may wear Feed pump, air Feed pump, nozzles Feed pump, rotary valve drive unit Parts that may Feed pipe Pump, feed line, Pump, feed line, corrode nozzles wheel Cost to replace worn Very low Higha) Higha) parts Time to replace worn Hours Daysb) Daysb) parts

a) Cost varies with product. b) Usually require spare pumps or spare atomizer assembly.

One major advantage of a strong oscillating hot gas jet from the tailpipe is that it can also be used to promote the dispersion of pasty and particulate materials. Figure 2.21 shows the flow diagrams of two prototype pulse combustion dryers designed by Novadyne Ltd (Canada). The pulse combustor uses a self-actuated flapper valve (see Fig. 2.7), and its operating frequency is about 70 Hz. The rated capacity of the combustor equal to 300 kW with a turndown ratio of 3 : 1 offers an evaporative capacity of about 2751 of kg wateh r. Drying tests proved that the thermal fiefciency of the pulse combustion fl ash dryer is similar to that of conventional flash dryers, but that the electrical power requirements were 40–50% lower as the motors used tended to be smaller and the need for certain material handling equipment (e.g., slinger) was eliminated (Buchkowski, 1999). The sample characteristics of a pulse combustion dryer for green hardwood sawdust and pulverized hog fuel are listed in Tab. 2.4. As some materials require longer residence times, the pneumatic duct (Fig. 2.21a) is replaced with a pressurized drying chamber connected directly to the tailpipe of the pulse combustor (Fig. 2.21b). Another possibility is to back-mix a fraction of the dry product from the cyclone with the highly wet feed in order to improve the flowability of the material entering the dryer. 2.5 Pulse Combustion in Drying 51

Fig. 2.21 (a) Pulse combustion pneumatic dryer and (b) pulse combustion dryer with pressurized chamber (courtesy of Novadyne Ltd, Hastings, ON, Canada).

Tab. 2.4 Test data for wood waste dried in a pulse combustion flash dryer (courtesy of Novadyne, Ltd, Hastings, ON, Canada).

Throughput (wet material) 590 kg h 1 Initial moisture content 45–55% wb Final moisture content 10–20% wb Temperature in the tailpipe 1040 C Exhaust gas temperature 104 C Combustion air flow rate 306 m3 h 1 Flue gases flow rate 1275 m3 h 1 Heat input 237 kW Evaporation rate 204 kg h 1 1 Unit heat consumption 4187 kJ kg H2O 1 Electrical energy consumption 0.019 kW kg H2O Back-mixing fraction 40% 52 2 Pulse Combustion Drying

Fig. 2.22 Basic configuration of the Pulse-Dryer (courtesy of Pulse Drying Systems Inc., Portland, OR, USA; www.pulsedrying.com).

Figure 2.22 illustrates, schematically, another pulse combustion dryer designed 1 by Pulse Drying Systems in Portland, Oregon, and marketed as . the Pulse-Dryer This dryer combines dispersion/drying in a tailpipe of the combustor and finish drying/separation in a primary cyclone. The air for combustion (and optionally quenching flue gases) is introduced into the system through an air column and is drawn into the combustor room and around the pulse combustor. The room can accommodate either one or more valveless pulse combustors operated at 250 Hz. Raw material is fed directly to the drying cone formed from a Venturi-shaped tailpipe of the combustor. Dry product and water vapor are blown from the drying cone into the primary collector, where the heavier particles are separated by gravity, while the vapor and fine particles continue to flow through the system to the secondary cyclone. The dry product falls out of the system through airlock valves, while the vapors are drawn through the spray tower before being vented to the atmosphere. The required moisture content is maintained by the air vent temperature, which controls the material feed rate. The Pulse-Dryer can handle products as diverse as fishmeal, brewer ’ s yeast, whey, various manures, sewage sludge, fruit and vegetable waste, kaolin clay and plating wastes. The particle size may vary from submicron to about 6 mm, and the moisture content may be as high as 99% wet basis. Interestingly, regardless of the feed rate of the raw material, the system removes an almost constant amount of water per hour from the feedstock, equal to approximately 1225 1 perkg hcombusto r. The single-module dryer shown in Fig. 2.22 can accommodate up to two pulse combustors, although the number of dryer modules can be increased as necessary to con figure a drying system with the required capacity. The average energy consumption of the Pulse-Dryer has been claimed to be about 3370 kJ 1 ofkg water evaporated, but it may be as low1 foras 2950 kJ kg finely dispersed materials with mostly surface water, such as kaolin clay, metal oxides or calcium carbonate (Duncan, 2012). Because the residence time in the References 53 drying cone is measured in milliseconds and the exit temperature of most products averages 52 C, the risk of thermal degradation is practically eliminated. At the same time, however, a much longer residence time in combination with relatively high temperatures in the primary cyclone (ranging from 80 to 120 C) creates conditions that reduce the content of pathogenic microorganisms, allowing the secure processing and disposal of sewage sludge and similar materials. Detailed information on these and other pulse combustion dryers is available in the book by Kudra and Mujumdar (2009).

Additional Notation Used in Chapter 2

f frequency s 1 a l Nu Nusselt number ( dp/ ) Pr Prandtl number (cm/l)  r m Re Reynolds number (udp / ) u average gas flow velocity m s 1 Sc Schmidt number (m/(dr)) b d Sh Sherwood number ( dp/ )

Abbreviations

CFD computational fluid dynamics db dry basis LDA laser Doppler anemometer psi pounds per square inch ppm parts per million PC pulse combustion PCD pulse combustion drying PDA particle dynamic analysis SPL sound pressure level SMW Sauter mean diameter VOC volatile organic compounds wb wet basis

References

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3 Superheated Steam Drying of Foods and Biomaterials Sakamon Devahastin and Arun S. Mujumdar

3.1 Introduction

Although most industrial dryers currently use hot air as the drying medium, it is well recognized that hot air drying leads to many adverse changes in terms of the dried product quality, including discoloration, shrinkage and deformation of shape, and losses of nutrients, among others. For this reason, a number of alternative drying technologies have been proposed and tested; these include such popular techniques as freeze drying, vacuum drying and microwave-assisted drying. Besides those drying technologies that have already been well adopted in industry, superheated steam drying (SSD) is another alternative that may be of interest, especially to the processors of foods and biomaterials. SSD has indeed been successfully applied, albeit mainly at a laboratory or pilot scale, to a wide array of products during the past decade (Mujumdar, 1990; Devahastin and Suvarnakuta, 2004; Mujumdar, 2007; Devahastin and Suvarnakuta, 2008). The basic principle of SSD is the use of superheated steam, both as the heating and drying medium, in a convective dryer. In principle, therefore, it should be possible to convert any convective hot-air dryer into a superheated steam dryer, although in practice this is rather complicated as a superheated steam dryer should be a closed system so as to avoid leakage of steam out of the dryer and infiltration of the air into the dryer. One major advantage of SSD is the ability to recover and reuse most of the energy supplied to the dryer, leading to substantial energy savings (as will be discussed below). SSD also possesses other important advantages, especially when applied to drying foods and biomaterials. One important benefit of SSD is that no (or a very small amount of) oxygen is present in the drying chamber, and this results in negligible oxidation reactions which, in most cases, are undesirable. Undesirable oxidation reactions that can be avoided when drying under the superheated steam environment are enzymatic browning reactions (Nimmol et al., 2007a; Thomkapanish et al., 2007), lipid oxidation, or the oxidative degradation of

Modern Drying Technology Volume 5: Process Intensification, First Edition. Edited by Evangelos Tsotsas and Arun S. Mujumdar. Ó 2014 Wiley-VCH Verlag GmbH & Co. KGaA. Published 2014 by Wiley-VCH Verlag GmbH & Co. KGaA. 58 3 Superheated Steam Drying of Foods and Biomaterials

vitamins (Suvarnakuta et al., 2005a; Methakhup et al., 2005; Kongsoontornkijkul et al., 2006; Hiranvarachat et al., 2008). In addition, it has been reported that SSD can effectively inactivate microorgan- isms, which can contaminate raw food materials. This is ascribed to the fact that SSD is generally performed at a higher temperature (above 100 C at atmospheric pressure). Moreover, the surface temperature of the drying material rapidly rises during the early period of drying due to steam condensation when a lower- temperature product is introduced into the SSD chamber. This rapid rise in the material’s temperature as the heat of condensation is released, in combination with the initially high water activity, results in an effective inactivation of microorgan- isms as compared to hot-air drying, which involves a much slower increase in the material’s temperature. It should be noted that the slower rise in the material’s temperature will allow microorganisms to proliferate or to produce toxins prior to their being inactivated (Chiewchan et al., 2007; Phungamngoen et al., 2011; Chiewchan, 2011). The use of superheated steam as the drying medium furthermore allows other necessary operations to be simultaneously conducted with the drying process. These include blanching (Iyota et al., 2001; Prachayawar- akorn et al., 2002; Namsanguan et al., 2004; Leeratanarak et al., 2006) and parboiling (Soponronnarit et al., 2006), as well as puffing or popping (Iyota et al., 2005). Another noted advantage of SSD is that the steam-dried product typically has a higher porosity than the same product dried in hot air. This is due to the boiling of water within the material during SSD; the boiling or vaporization of water leads to a significant volume expansion due to a rapid pressure build-up if the steam generated cannot diffuse out fast enough and, as a result, expands the microstructure (i.e., a higher porosity) of the dried material. This allows the dried product to be rapidly reconstituted upon immersion in water, which is a highly desirable characteristic of any “instant” food.

3.2 Principle of Superheated Steam Drying (SSD)

Figure 3.1 illustrates the principle of SSD, which is started by feeding saturated steam produced by a boiler to a heater (steam superheater) so as to convert the saturated steam into superheated steam. The superheated steam is then fed into a drying chamber where the steam transfers heat to the drying material and picks up the evaporated moisture from the drying material. This results in a lower temperature and a specific enthalpy of the exhaust steam. The exhaust steam can (and should) be reheated and reused in a closed loop. The extra steam generated in the dryer should also be withdrawn from the system and used elsewhere. Provided that the exhaust steam can be reused (or condensed and its energy used elsewhere), the “net” energy consumption of the drying system would be much reduced, as the latent heat of vaporization is not charged to the dryer operation. In such a case, the required energy for drying is simply the sensible heat that is required to reheat the saturated steam into superheated steam; the energy required to produce saturated 3.2 Principle of Superheated Steam Drying (SSD) 59

Fig. 3.1 Basic superheated steam drying system. Modified from Devahastin and Suvarnakuta (2008).

steam needs to be counted only once. In practice, however, it may be impossible to reuse all of the exhaust steam but, based on the present authors’ experience, a 60– 70% recovery is possible. This nevertheless represents a significant energy saving compared to a hot-air drying system. If the liquid to be removed from a drying material is water, there would be no resistance to moisture transfer from the material surface into superheated steam since the drying medium and the liquid to be removed are the same species. For this reason, it is possible to calculate the rate of drying in a constant drying rate period only from the steam-to-material heat-transfer data; the rate of moisture evaporation would simply vary approximately linearly with the rate of heat transfer to the drying material. If it is assumed that there are no heat losses from the drying chamber, and that drying is purely by convection (no such other modes of heat transfer as conduction or radiation), the rate of drying in a constant drying rate period is given by: q hTðÞ T m_ ¼ ¼ steam surf ð3:1Þ Dhv Dhv;surf where m_ is the rate of drying (or rate of moisture evaporation from the material fi surface), h is the heat transfer coef cient, Tsteam is the superheated steam temperature, Tsurf is the material surface temperature, which is equal to the saturation temperature of water at the dryer operating pressure, and Dhv;surf is the latent heat of vaporization at the material surface temperature and dryer operating pressure. Note that there are several simplifying assumptions involved in Eq. 3.1. When calculating the rate of drying in a falling rate period, a more detailed mathematical model is needed (Kittiworrawatt and Devahastin, 2009). Another feature of SSD is the so-called inversion phenomenon. This implies a higher rate of drying in a constant drying rate period when using superheated 60 3 Superheated Steam Drying of Foods and Biomaterials

steam as the drying medium than the rate of drying in the corresponding period when using hot air as the drying medium. This phenomenon takes place beyond the so-called inversion temperature; at this temperature the rates of drying in a constant drying rate period when using superheated steam and hot air as the drying media are the same. The reason for the inversion phenomenon can be seen

from Eq. 3.1. In the case of hot-air drying, Tsurf in Eq. 3.1 would be equal to the wet- bulb temperature of the drying air. At lower drying temperatures, the temperature difference in the case of hot-air drying is thus higher than in the case of SSD. This is because the wet-bulb temperature is generally significantly lower than the hot-air drying (or the dry-bulb) temperature, resulting in the relatively higher rate of heat transfer, and hence higher rate of drying. However, as the drying temperature increases the temperature difference in the case of SSD would become much higher, while the temperature difference in the case of hot-air drying would be increased only moderately. This results in the rate of SSD being equal to the rate of hot-air drying at a particular temperature – that is, the inversion temperature (see Fig. 3.2). For more information on the inversion phenomenon and inversion temperatures of different systems, the reader is referred to Sheikholeslami and Watkinson (1992), Schwartze and Brocker (2002) and Prachayawarakorn et al. (2002), among others. It should be noted that this inversion phenomenon is not very important in deciding whether to use SSD in preference to hot-air drying. Other considerations, such as better product quality, safety and energy savings, are more important factors affecting the selection. Obviously, the inversion phenomenon can be defined only for a constant drying rate period. Nevertheless, several research groups have attempted to extend this definition also to a falling rate period. Prachayawarakorn et al. (2002), for example, have reported the inversion temperature for the whole drying process (combined

Fig. 3.2 Illustration of the inversion phenomenon. Modified from Devahastin and Suvarnakuta (2008). 3.3 Atmospheric-Pressure Superheated Steam Drying 61 constant and falling rate periods). Care should therefore be taken to consider if the reported inversion temperature belongs only to a constant drying rate period or to the combined constant and falling rate periods. As expected, the inversion temperature calculated only from the constant drying rate and that calculated from the combined drying rate are not equal (Suvarnakuta et al., 2005b), since the latter depends on the material properties as well as the drying conditions.

3.3 Atmospheric-Pressure Superheated Steam Drying

Superheated steam drying at atmospheric (or near atmospheric) pressure is the process that has been most extensively studied compared to SSD at other pressure ranges – that is, at a pressure higher or lower than the atmospheric pressure. In this section, some examples are given to explain the consequences of SSD, from both the drying (heat/mass transfer) and product quality changes points of view. This overview will be limited to foods and biomaterials, and hence will exclude products such as wood, paper, ceramics, sludge, coal, and peat that can also be dried in a superheated steam environment. Li et al. (1998, 1999) investigated the use of superheated steam to dry tortilla chips in an impingement dryer. When they first studied the effects of superheated steam temperature and flow rate (or velocity) on the drying kinetics and quality of dried tortilla chips in terms of crispness, starch gelatinization and microstructure, the steam temperature was noted to have a more dominant effect on the drying kinetics than the steam flow rate, since the drying took place in a falling rate period. In terms of the dried product quality, the use of a higher superheated steam temperature led to a more complete starch gelatinization; the chips prepared at a higher superheated steam temperature also reconstituted better than those prepared at a lower superheated steam temperature. This was due to the larger pores that developed during the higher-temperature drying due to boiling of internal moisture (see Fig. 3.3, which illustrates the effect of steam flow rate in terms of the convective heat transfer coefficient); the pore formation was the result of the boiling of moisture mentioned earlier. The observed results were consistent with those of Caixeta et al. (2002), who conducted experiments with potato chips. However, the latter group reported that the drying of potato chips occurred only in a falling rate period if hot air was used as the drying medium. This was opposite to the use of superheated steam as the drying medium, where drying occurred both in constant and falling rate periods; SSD therefore represented a faster drying technique. In terms of the dried product quality, potato chips dried at higher superheated steam temperatures suffered less shrinkage than those dried at lower superheated steam temperatures; nevertheless, the chips dried at higher super- heated steam temperatures suffered more color changes and exhibited a lower vitamin C retention. When compared with hot-air dried chips, the steam-dried chips suffered more extensive shrinkage, most likely because the hot-air-dried chips developed casehardened surfaces that helped inhibit shrinkage. Nevertheless, 62 3 Superheated Steam Drying of Foods and Biomaterials

T = 115 °C; h = 110 W m-2 K-1 T = 145 °C; h = 110 W m-2 K-1

T = 115 °C; h = 160 W m-2 K-1 T = 145 °C; h = 160 W m-2 K-1

Fig. 3.3 Microstructural images of tortilla chips dried at different conditions. Reproduced with permission from Li et al. (1999).

the steam-dried chips were less hard and crispier than the air-dried chips at the same drying conditions. Steam-dried chips also exhibited a higher vitamin C retention than the air-dried chips, especially at a lower drying temperature, as shown in Tab. 3.1. Iyota et al. (2001) dried potato chips using both hot air and superheated steam as the drying media. In terms of the drying kinetics, it can be seen from Fig. 3.4 that the use of hot air (which had an absolute humidity of 0.02 kg kg 1) led to a higher drying rate than when using superheated steam as the drying medium at the same drying temperature of 170 C. However, when considering both drying processes at 240 C, the rates of drying were seen to be almost identical, with the two drying curves almost collapsing into one curve. Thus, it is possible to conclude that the inversion temperature for the whole drying process (see Section 3.2 for explana- tion) was around 240 C. Steam condensation on the material surface during an early period of SSD can also be clearly seen from the slight increase in the moisture content in such a period. When examining the near-surface cross-sections of the chips with scanning electron microscopy (SEM), it was noted that many starch granules still existed in the hot-air-dried chips. However, an opposite effect was noted for the steam-dried chips, which contained practically no intact starch granules (see Fig. 3.5). These

Tab. 3.1 Vitamin C contents of potato chips dried at different conditions. Modified from Caixeta et al. (2002).

Drying temperature (C) Vitamin C content (mg per 100 g d.b.)

Superheated steam drying Hot air drying

130 25.48 1.25 21.43 0.57 145 20.48 1.13 19.06 2.71 3.3 Atmospheric-Pressure Superheated Steam Drying 63

Fig. 3.4 Drying curves of potato chips at different conditions. Modified from Iyota et al. (2001).

findings implied a more complete gelatinization of starch in the steam-dried chips. The microstructures of the hot-air-dried and steam-dried chips were also quite different; the steam-dried chips possessed a more porous structure than the hot-air- dried chips, and a casehardened skin was clearly visible in the hot-air-dried chips. The occurrence of a casehardened skin, combined with a lower porosity, led to the hot-air-dried chips having an inferior texture and rehydration characteristics. Additional results and an explanation related to this important point are provided in a later section of this chapter. A number of other food materials have also been dried in superheated steam. Tang et al. (2000), for example, dried sugar beet pulp using SSD and compared the results with those obtained from hot-air drying experiments. The medium temperature was noted to have a more significant effect on the required drying time in the case of SSD than in the case of hot-air drying, most likely because the thermal properties of superheated steam are more sensitive to the temperature than are those of hot air. The mass of the sample slightly increased during the early period of drying, again as a result of steam condensation, and condensation occurred more at a lower superheated steam temperature, as expected. Another interesting result (see Fig. 3.6) was that the rate of drying was unaffected by the hot air velocity, but was significantly affected by the superheated steam velocity at sample moisture contents higher than X ¼ 0.5. This indicated that the critical moisture content of the sugar beet pulp was lower when being dried by superheated steam; the gas-side-controlled drying rate period was therefore longer in the case of SSD. Both, the steam-dried and hot-air dried samples nevertheless exhibited similar sorption behaviors. 64 3 Superheated Steam Drying of Foods and Biomaterials

Fig. 3.5 Microstructural images of potato chips dried at different conditions. Modified from Iyota et al. (2001).

Fig. 3.6 Hot air drying (HAD) and superheated steam drying (SSD) rates of sugar beet pulp. Reproduced with permission from Tang et al. (2000). 3.3 Atmospheric-Pressure Superheated Steam Drying 65

Prachayawarakorn et al. (2002) studied the drying kinetics of shrimp undergoing hot-air drying and SSD within the temperature ranges of 70 to 140 C and 120 to 180 C, respectively. They also proposed a simple mathematical model that can be used to simulate the changes of the shrimp moisture content in a falling rate period. It was shown that the effective diffusion coefficient of moisture within the shrimp was different when using different drying media. The difference can be seen from the following Arrhenius-based relationships between the effective fi 2 1 moisture diffusion coef cient Deff (in m s ) and the medium temperature T (in K):  32714:9 SSD : Deff ¼ 0:00253 exp ð3:2Þ RT~  25578:0 Hot-air drying : Deff ¼ 0:00253 exp ð3:3Þ RT~ where R~ is the universal gas constant in J kmol 1 K 1. More appropriately, and theoretically correctly, however, the material temperature should be used instead of the medium temperature. If the effective moisture diffusion coefficients are calculated from Eqs 3.2 and 3.3, it can be noted that the coefficient in the case of SSD would be higher when the medium temperature is higher than 140 C. This corresponds to the experimental drying rate data shown in Fig. 3.7; it can be seen that the inversion temperature of the whole drying process was approximately 140 C.

Fig. 3.7 Hot air drying (HAD) and superheated steam drying (SSD) rates of shrimp. Reproduced with permission from Prachayawarakorn et al. (2002). 66 3 Superheated Steam Drying of Foods and Biomaterials

Tab. 3.2 Processes, conditions and shrinkage for multistage shrimp drying. Modified from Namsanguan et al. (2004).

Case Process/Condition Shrinkage (%)

1 Only SSD at 140 C 37.5 7.5b 2 SSD at 140 C followed by HPD at 50 C (IMC 30% w.b.) 21.5 7.4b 3 SSD at 140 C followed by HPD at 50 C (IMC 40% w.b.) 31.2 5.7a,b 4 SSD at 140 C followed by HAD at 50 C (IMC 30% w.b.) 35.5 7.2b 5 SSD at 140 C followed by HAD at 50 C (IMC 40% w.b.) 48.1 8.9b

The same superscript letters indicate that the mean values of shrinkage are not significantly different when considered at a confidence level of 95% (p > 0.05). HAD: Hot air drying; HPD: Heat pump drying; IMC: Intermediate moisture content.

If there is a desire to further improve the quality of a dried product, then a multistage drying process may be employed. Such a process can be realized by conducting SSD during the first stage, followed by a second-stage drying process at a lower temperature to avoid exposing the drying material to the high-temperature drying medium (viz. superheated steam) for an extended period of time. Namsanguan et al. (2004) validated the aforementioned concept by drying shrimp in a multistage fashion. The process started with SSD at 140 C, followedbyheatpumpdrying(HPD)at50C. The results were compared with those obtained through first-stage SSD at 140 C and then hot-air drying (HAD) at 50 C; single-stage SSD at 140 C was also conducted, and the results were compared. These investigators also studied the effect of varying the intermedi- ate moisture content (IMC), which is the moisture content of shrimp at the end of the first drying stage (in this case SSD), to determine how the drying time of the different drying processes would affect the overall drying kinetics and the quality of the dried product. Indeed, a higher IMC implied that the time spent by the material during the first drying stage is shorter. All of the tested conditions are listed in Tab. 3.2. Figure 3.8a, which illustrates the drying kinetics of shrimp undergoing different drying process combinations, clearlyshowsthatthemoisturecontent of shrimp decreased most rapidly if only SSD was applied (Case 1). This is because the shrimp was exposed to the high-temperature superheated steam throughout the whole period of drying. In the cases of multistage drying, however, a higher IMC resulted in a longer overall drying time as the period with the higher drying rate (SSD) was shorter, as mentioned earlier. Never- theless, it is interesting to see from Fig. 3.8b that SSD/HPD required as much as a 1.4 to 1.5-fold shorter drying time than SSD/HAD at the same IMC. This point is ascribed to the fact that HPD involves drying at a much lower air relative humidity as compared to HAD. In terms of the dried product quality, it was noted that the shrimp which had undergone multistage drying (especially SSD/HPD) suffered significantly less shrinkage than the shrimp dried only by SSD (see Tab. 3.2). This was due to the fact that superheated steam-dried shrimp was exposed to the high-temperature 3.3 Atmospheric-Pressure Superheated Steam Drying 67

Fig. 3.8 (a) Drying curves and (b) drying times of shrimp at different conditions. Modified from Namsanguan et al. (2004). drying medium for an extended period of time, leading to a collapse of the porous structure that had been formed as a result of water vaporization during an early period of SSD; in the case of SSD, the shrimp temperature eventually reached a value which was much higher than the glass transition temperature, 68 3 Superheated Steam Drying of Foods and Biomaterials

Fig. 3.9 Microstructural images of shrimp dried at different conditions. (a) Only SSD at 140 C; (b)SSDat140 CfollowedbyHPD(IMC40%w.b.);(c)SSDat140 CfollowedbyHPD(IMC30%w.b.). Reproduced with permission from Namsanguan et al. (2004).

and this resulted in a rather severe phase transition and collapse. On the other hand, if the partially superheated steam-dried shrimp, which still possessed an intact high-porosity structure, was subsequently transferred into a lower- temperature drying environment, the structure of shrimp would be much better preserved, with less shrinkage. This hypothesis was indeed verified by the microstructural images of shrimp shown in Fig. 3.9, which clearly show that the shrimp which had undergone SSD/HPD at a lower IMC exhibited a higher- porosity structure than the shrimp that had been dried only by SSD. However, it can be seen from the data in Tab. 3.2 that the shrimp which undergone SSD/ HAD had a similar (or even higher) level of shrinkage than those dried only by SSD. This may be because SSD/HAD was a relatively slow process compared to SSD and SSD/HPD, especially when the IMC was higher. A longer drying process is indeed known to lead to a greater degree of structural collapse and shrinkage of the drying material. In terms of color, shrimp subjected to SSD/HPD was seen to possess a redder color than shrimp dried only by SSD. The shrimp dried by SSD/HPD was also softer; the reason for these observations being similar to those applied earlier to explain the collapse and shrinkage phenomena. One major advantage of using superheated steam to dry shrimp is that the boiling stage, which normally needs to be conducted prior to drying, can be 3.4 Low-Pressure Superheated Steam Drying (LPSSD) 69 eliminated, and the shrimp can simply be immersed in saline solution and introduced to SSD. Boiling and drying would then occur simultaneously. This represents an alternative approach to reducing the time and energy consumption of the entire dried shrimp production process. Moreover, dried shrimp obtained via such a process also taste better as the flavoring ingredients are not lost with the boiling solution.

3.4 Low-Pressure Superheated Steam Drying (LPSSD)

Since atmospheric-pressure (or near atmospheric-pressure) SSD involves the use of a high-temperature drying medium (i.e., superheated steam at a temperature higher than 100 C), this type of process may not be appropriate for heat-sensitive materials such as fruits, vegetables, herbs and spices, and many other biomaterials. For this reason, an alternative drying technique has been developed to combine the advantages of SSD with an ability to conduct drying at a lower temperature. The proposed alternative is based on the fact that water vaporizes at a lower temperature if the pressure of the system is lower; superheated steam can then be produced at a lower temperature under such a lower pressure. The so-called low-pressure superheated steam drying (LPSSD) process has been tested with a wide array of foods and biomaterials, with the first comprehensive test covering both the drying kinetics and quality of the dried product (carrots) having been conducted by Devahastin et al. (2004); a schematic diagram of the employed LPSSD set-up is shown in Fig. 3.10. This type of set- up can be attached to a boiler, which is available in a typical food-processing plant. However, as a typical boiler generates saturated steam at a pressure of around 6–8 bar, the steam cannot be directly introduced into the drying chamber. A reservoir, along with a pressure regulator, is installed to maintain the saturated steam at around 1.5–2 bar prior to feeding it to the drying chamber. Once the saturated steam flows into the low-pressure drying chamber, the steam would spontaneously become superheated steam as the temperature of the entering steam is already higher than the saturation temperature of water at the pressure of the drying chamber. Nevertheless, to allow a more precise control of the temperature within the drying chamber, an electric heating coil is installed to help regulate the superheated steam temperature. At an industrial- scale setting, this electric heating coil may be replaced by a coil through which hot oil or high-pressure steam can run. When carrots were subject to drying at 60–80 C over a pressure range of 7–13 kPa, they were found to experience a slight increase in moisture content during the early period of drying, which is a typical characteristic of any superheated steam drying process. The initial condensation decreased on reducing the pressure of the drying chamber, as the reduced boiling point at a lower pressure (at a constant superheated steam temperature) led to a 70 3 Superheated Steam Drying of Foods and Biomaterials

Fig. 3.10 Low-pressure superheated steam (12) Temperature sensor and recorder; drying (LPSSD) set-up. (1) Boiler; (2) Steam (13) Vacuum break-up valve; (14) Insulator; valve; (3) Steam reservoir; (4) Pressure gage; (15) Load cell and mass recorder; (16) Vacuum (5) Steam trap; (6) Steam regulator; (7) LPSSD pump; (17) PC and data acquisition system. chamber; (8) Steam inlet; (9) Electric fan; Reproduced with permission from Devahastin (10) Sample holder; (11) Electric heater; et al. (2004).

higher degree of superheating (i.e., the difference between the superheated steam temperature and saturation temperature of water at a specific pressure). In addition, the drying temperature was noted to have a more significant effect on the drying kinetics of the product than did the pressure, notably because the temperature has a more dominant effect on the steam thermophysical properties than the pressure. Nevertheless, the drying rate of carrots increased as the pressure of the drying chamber decreased (at a constant superheated steam temperature), as the boiling point of water decreasedatalowerpressureandthisresultedinafasterremovalofwater from the carrots. If a comparison is made between LPSSD and vacuum drying (VD), the time needed to dry a material using VD would be less than that required for LPSSD at the same temperature and pressure. This difference can be ascribed to the fact that the material receives more energy from the heater in the case of VD because the heater is the sole energy source for drying; thus, the material absorbs more radiated energy from the heater and exhibits a higher temperature and, hence, a higher drying rate. Nevertheless, the difference in the required drying time would be smaller at a higher drying temperature, which is indeed a characteristic of the inversion phenomenon. In order to preserve the quality of the dried product, however, it is not always recommended to conduct LPSSD at the inversion temperature or higher; it should be noted that the inversion temperature in the case of LPSSD may be as high as 90 C (Suvarnakuta et al., 2005b). Figure 3.11 illustrates the drying curves and temperature evolution patterns of a material (carrots) undergoing LPSSD. These data show clearly that the drying temperature and pressure both have significant effects on the moisture and 3.4 Low-Pressure Superheated Steam Drying (LPSSD) 71

Fig. 3.11 Evolutions of moisture content and temperature of carrots during LPSSD. (a) 60 C, 7 kPa; (b) 70 C, 7 kPa; (c) 80 C, 7 kPa; (d) 80 C, 13 kPa. ~: moisture content; : steam temperature; : carrot temperature. Reproduced with permission from Devahastin et al. (2004). temperature evolution patterns. In general, the material temperature would increase rapidly from the initial temperature to the saturation temperature of water corresponding to the pressure of the drying chamber, and would then remain at the saturation temperature until the end of a constant drying rate period. The material temperature would subsequently increase and approach the superheated steam temperature towards the end of drying. One important observation that can be made upon examining Fig. 3.11 is that when the superheated steam temperature increases (at a constant pressure), the period where the material temperature would remain constant (i.e., the constant drying rate period) would become shorter. In some cases where the superheated steam temperature is higher and the pressure of the drying chamber is lower (e.g., at a temperature of 80 C and a pressure of 7 kPa; see Fig. 3.11c), the material temperature would continuously increase and approach the superheated steam temperature without having a constant- temperature period. In contrast, if the superheated steam temperature is lower and the pressure of the drying chamber is higher, then the period of constant material temperature would be longer (see Fig. 3.11d). These findings confirm 72 3 Superheated Steam Drying of Foods and Biomaterials

that a constant drying rate period – or, in other words, the critical moisture content – is dependent on not only the type of material but also the drying conditions. Figure 3.12 illustrates sample drying curves and temperature evolution patterns of a material (again carrots) undergoing VD at the same drying conditions as those shown in Fig. 3.11. It can be seen that the moisture and temperature evolution patterns (or heat and mass transfer behavior) of the material are quite different from those of the sample undergoing LPSSD. In the case of VD, the material temperature continuously increases toward the drying temperature without having a constant-temperature period. However, the vacuum-drying rates are seen to be higher than the LPSSD rates at the same operating conditions, as mentioned earlier. In terms of dried product quality, the volume and density of the carrots changed in a similar manner upon both VD and LPSSD. However, although both vacuum- dried and low-pressure superheated steam-dried carrots exhibited similar volu- metric shrinkage (calculated from the change in the volume of a sample compared

Fig.3.12 Evolutionsofmoisturecontentandtemperatureofcarrotsduringvacuumdrying.(a)60 C, 7kPa;(b)70C, 7 kPa; (c) 80 C, 7 kPa; (d) 80 C, 13 kPa. ~, moisture content;, drying temperature; , carrot temperature. Reproduced with permission from Devahastin et al. (2004). 3.4 Low-Pressure Superheated Steam Drying (LPSSD) 73

Fig. 3.13 Photographic images of dried and reconstituted (rehydrated) carrots. to the initial volume of the sample prior to drying), the deformation patterns of the two sets of samples were quite different. Carrots which had undergone VD suffered a much more nonuniform shrinkage than those that had undergone LPSSD. This difference may be because VD led to a higher drying rate and, in turn, causing casehardening of the carrot surface before the inner part of the carrot could be dried. In other words, the inner carrot dried much more slowly than the surface, and this naturally led to a nonuniform shrinkage (see Fig. 3.13). It can be seen from the microstructural images shown in Fig. 3.14 that carrots dried by LPSSD exhibited a more uniform microstructure than those dried by VD. Notably, because the superheated steam-dried carrots deformed more uniformly, they were more easily reconstituted when immersed in water – a very desirable characteristic sought by the instant food industry. In addition to studies on the changes of various physical characteristics of materials undergoing LPSSD, several investigations have been made into the use of this drying technique to preserve various heat-sensitive bioactive compounds. For

Fig. 3.14 Microstructural images of carrots after (a) LPSSD and (b) vacuum drying. Reproduced with permission from Devahastin et al. (2004). 74 3 Superheated Steam Drying of Foods and Biomaterials

example, Methakhup et al. (2005) prepared dried Indian gooseberry (Phyllanthus emblica L.), which can be used as an ingredient of a herbal tea, by using both VD and LPSSD. Although LPSSD was noted to require a longer drying time, the technique resulted in a higher retention of vitamin C (ascorbic acid) in the dried gooseberry than did VD (see Tab. 3.3). In addition, the gooseberry dried by LPSSD exhibited much less of a color change than did gooseberries treated with VD. It can also be seen from the data in Tab. 3.3 that the change in the VD temperature significantly affected the ascorbic acid content and the color of the dried gooseberry, whereas the superheated steam drying temperature (and pressure) had no significant effects on the quality parameters of the dried gooseberry. These findings may help confirm the negligible oxygen environment of LPSSD, which led to a negligible oxidation of ascorbic acid and other phenolic compounds in the gooseberry, and thus resulted in smaller losses of ascorbic acid and less color change. Rather, the loss of ascorbic acid and color changes after LPSSD were considered due to thermal degradation and other reactions that do not involve oxygen, including nonenzymatic browning reactions (e.g., the Maillard reaction, which occurs between amino acids and reducing sugars). In addition to the retention of vitamin C in carrots, both VD and LPSSD can also significantly reduce losses of b-carotene when compared to hot air drying, as demonstrated by Suvarnakuta et al. (2005a) (see Tab. 3.4). This effect is, again, due to the lack of oxygen during VD and, in particular, during LPSSD. The retention of b-carotene in carrots undergoing VD and LPSSD did not differ greatly, unlike that of ascorbic acid, because the oxygen content plays a more significant role in the loss of the latter than of the former. Thus, it can be concluded that LPSSD has clear advantages over other competing drying techniques when applied to heat-sensitive and, in particular, oxygen-sensitive materials. From the consideration of food safety, which involves the inactivation or growth- suppression of microorganisms, LPSSD has been noted to be more effective than other commonly used drying techniques at the same drying temperature. For example, Phungamngoen et al. (2011) tested two different drying techniques that have been commonly applied to drying foods and biomaterials – namely hot air drying and VD – and compared these with LPSSD over a temperature range of 50 to 70 C. The three drying techniques were evaluated for their effectiveness in decontaminating a pathogenic bacterium (Salmonella Anatum), which was inoculated onto the surfaces of cabbage leaves. The results, shown in terms of the survival curves of S. Anatum (see Fig. 3.15), showed that LPSSD reduced the bacterial numbers faster than did hot air drying and VD, especially at 60 and 70 C, despite requiring 30% and 14% longer times than VD to reduce the cabbage moisture content to the desired final value at 60 and 70 C, respectively. This is because the temperature of the cabbage rose rapidly in the case of LPSSD. The initial condensation, which occurred only in the case of LPSSD, also helped increase (or at least maintain) the activity of water on the surfaces of the cabbage. As noted earlier, the inactivation of microorganisms is more effective under a high water activity condition (Chiewchan, 2011), and this renders LPSSD more effective for microbial inactivation than the alternative drying techniques that were tested.

76 3 Superheated Steam Drying of Foods and Biomaterials

Tab. 3.4 b-carotene contents of fresh carrot and carrots dried at different conditions. Modified from Suvarnakuta et al. (2005a).

Sample b-carotene content (mg per 100 g 1) Retention (%)

Dry basis Wet basis

Fresh carrot 4.86 0.65 51.11 6.89 — Carrots dried by LPSSD 60 C 54.92 43.94 83 2a 70 C 50.32 41.24 76 4b 80 C 38.81 31.55 74 2b Carrots dried by vacuum drying 60 C 26.95 24.28 74 3c 70 C 25.81 24.05 71 3c,d 80 C 31.61 27.85 68 3d Carrots dried by hot air drying 60 C 22.56 21.46 62 2e 70 C 35.66 23.30 62 3e 80 C 36.26 22.05 58 3e

Same letters in the same column indicate that the mean values are not significantly different when considered at a confidence level of 95% (p > 0.05).

Figure 3.16a shows that S. Anatum contaminated on the surface of raw cabbages exhibited an intact rod shape, and the bacterial cells were distributed quite uniformly on the leaves. Upon drying, however, the cells suffered significant mor- phological changes, especially in the cases of VD and LPSSD (see Figs. 3.16c and d).

3.5 Application of LPSSD to Improve the Quality of Foods and Biomaterials

It is generally recognized that foods and biomaterials suffer quality degradation, in one way or another, upon drying. However, if an appropriate drying technique, possibly combined with appropriate drying conditions, can be employed then it might be possible to improve the quality or selected properties of foods and biomaterials, and perhaps even make the dried products superior to the fresh materials. Some examples of the use of LPSSD, in combination with other assisting drying processes or heating methods, are briefly reviewed in the present section. Emphasis is placed on the production of fat-free snacks that currently attract attention from health-conscious consumers as an alternative to deep-fried snacks. Highly crisp, fat-free snacks may be produced through the combined use of LPSSD and far-infrared radiation (FIR). For example, Nimmol et al. (2007a, 2007b) applied FIR-assisted LPSSD (LPSSD-FIR) to produce dried banana chips, while FIR-assisted vacuum drying (VACUUM-FIR) and LPSSD were also evaluated over 3.5 Application of LPSSD to Improve the Quality of Foods and Biomaterials 77

Fig. 3.15 Survival curves of Salmonella Anatum the number of bacteria at t ¼ 0). (a) Hot air contaminated on cabbage surfaces undergoing drying; (b) Vacuum drying;(C) LPSSD.&:50C; different drying processes and conditions (log S ^:60C; Î:70C. Reproduced with isthelogofthenumberofbacteriaatanytimetto permission from Phungamngoen et al. (2011).

the drying temperature range of 70–90 C at pressures of 7 and 10 kPa. The banana temperature (see Figs 3.17 and 3.18) was seen to slightly decrease from the initial temperature during an early period of drying. This was due to a reduction in the pressure of the drying chamber, leading to a flash evaporation of the surface moisture and hence a cooling of the banana. Subsequently, the banana temperature in the case of LPSSD-FIR rapidly increased and approached (but did not equal) the saturation temperature of water at the dryer operating pressure. The temperature evolution pattern in this case was quite different from that of LPSSD, where the material temperature increased to the saturation temperature of water at 78 3 Superheated Steam Drying of Foods and Biomaterials

Fig. 3.16 Scanning electron microscopic LPSSDfor1 h.Thedryingtemperaturewas60 C images of Salmonella on cabbage surfaces. in all cases. Reproduced with permission from (a) Prior to drying (fresh cabbage); (b) After hot Phungamngoen et al. (2011). air drying for 6 h; (c) After VD for 1 h; (d) After

the dryer operating pressure. This difference was due to the use of FIR as an extra source of energy for drying; drying is no longer purely convective when the FIR is present. When a constant drying rate period had elapsed, the banana temperature was noted to increase to a value that was higher than the superheated steam temperature; this was again due to the existence of FIR. Nevertheless, the banana temperature did not rise further as the fruit possessed much less moisture at this point and hence was unable to absorb any more radiated energy. It was noted further that, at the same drying temperature, the banana temperature towards the end of the drying process in the case of LPSSD-FIR (Fig. 3.17) was higher than that in the case of VACUUM-FIR (Fig. 3.18). This was due to the fact that during LPSSD-FIR, the radiation intensity at the position of a temperature sensor used to measure the medium temperature (which was in turn used to control the infrared radiator) was lower than that in the case of VACUUM- 3.5 Application of LPSSD to Improve the Quality of Foods and Biomaterials 79

Fig. 3.17 Evolutions of moisture content and moisture content; : drying temperature; : temperature of bananas during LPSSD-FIR at banana temperature. Reproduced with different conditions. (a) 80 C, 7 kPa; (b) 80 C, permission from Nimmol et al. (2007b). 10 kPa; (c) 90 C, 7 kPa; (d) 90 C, 10 kPa. Î:

FIR. This situation occurred because part of the radiated energy had been absorbed by the superheated steam, which possesses a higher ability to absorb FIR than does the air. For this reason, in the case of LPSSD-FIR, the radiator must be switched on more frequently in order to maintain the required radiation level, leading to a higher surface temperature of the banana than in the case of VACUUM-FIR. In terms of dried product quality, banana chips produced by LPSSD-FIR exhibited a darker color than those produced by VACUUM-FIR in all cases. This was because the banana temperature during LPSSD-FIR was higher than that during VACUUM-FIR, as mentioned earlier. The higher temperature expectedly led to accelerated rates of browning reactions, and hence to darker colors. It should be noted that the dominant browning reactions were nonenzymatic in nature, as the amount of oxygen in the drying systems was very low (almost negligible in the case of LPSSD-FIR). In addition, when comparing the processes with and without FIR, it was noted that the color of banana chips which had undergone the processes without FIR suffered less changes. This 80 3 Superheated Steam Drying of Foods and Biomaterials

Fig. 3.18 Evolutions of moisture content and moisture content; : drying temperature; : temperature of bananas during VACUUM-FIR at banana temperature. Reproduced with different conditions. (a) 70 C, 7 kPa; (b) 80 C, permission from Nimmol et al. (2007b). 10 kPa; (c) 90 C, 7 kPa; (d) 90 C, 10 kPa. Î:

was again related to the changes in the material temperature during drying (Nimmol et al., 2007a; Thomkapanish et al., 2007). Interestingly, the radial shrinkage of bananas subjected to lower-temperature drying (e.g., at 70 and 80 C) was noted to be less than that of bananas undergoing drying at a higher temperature. This was ascribed to the fact that casehardened skins had been formed during lower-temperature drying, and these had helped to retard the radial shrinkage. Both LPSSD-FIR and VACUUM-FIR were also noted to lead to a more extensive formation of the casehardened skins than LPSSD alone, again due most likely to the different temperature evolution patterns of the processes with and without FIR (as mentioned earlier). In terms of rehydration behavior, which is an important characteristic of any instant food, the banana chips produced by drying at a higher temperature were better able to absorb water than those produced at a lower temperature. This was because high-temperature drying had led to chips with a higher porosity as a result of a more intensive vaporization of the internal water (as illustrated comparatively in Fig. 3.19). LPSSD-FIR was noted to produce chips with a better rehydration ability than VACUUM- FIR for reasons related to the changes in the material temperature and the internal 3.5 Application of LPSSD to Improve the Quality of Foods and Biomaterials 81

Fig.3.19 Microstructuralimagesofbananachipsdriedatdifferentconditions.(a)LPSSD-FIRat80 C and 7 kPa; (b) VACUUM-FIR at 80 C and 7 kPa; (c) LPSSD-FIR at 90 C and 7 kPa; (d) VACUUM-FIR at 90 C and 7 kPa. Reproduced with permission from Nimmol et al. (2007a). vaporization of water (as described in Section 3.4). However, when drying was performed at a higher temperature of 90 C, the chips prepared by both processes exhibited similar rehydration behaviors. This most likely occurred because the levels of internal water vaporization of both groups of samples were no longer different, and this led to insignificantly different porosities of the chips (Nimmol et al., 2007a). The results of textural evaluations of banana chips prepared under different conditions are listed in Tab. 3.5. These results are expressed in terms of the maximum force in a force-deformation curve, which is a representative of the hardness of the samples. The results are also expressed in terms of the number of peaks in the force-deformation curve, which is a representative of the crispness of the samples. It can be seen that chips prepared by VACUUM-FIR were harder than those prepared by LPSSD-FIR. The difference was due to variations in the microstructure of the chips, as illustrated in Fig. 3.19. The drying temperature and pressure nevertheless had no significant effect on the hardness of the chips. In terms of crispness, LPSSD-FIR resulted in chips with higher numbers of peaks (and hence were crispier) than did VACUUM-FIR, especially when drying was conducted at a lower temperature, such as 80 C. This was again due to the more extensive porous structure of the chips produced by LPSSD-FIR. When combining LPSSD with FIR, it is clear that the hardness of the chips significantly decreased, whereas the crispness was significantly increased 82 3 Superheated Steam Drying of Foods and Biomaterials

Tab. 3.5 Texture of banana chips dried at different conditions. Modified from Nimmol et al. (2007a); LPSSD data from Thomkapanish et al. (2007).

Drying Temperature (C) Pressure (kPa) Maximum Number method force (N) of peaks

LPSSD-FIR 70 7 N/A N/A 10 N/A N/A 80 7 17.09 3.15a 37 3d 10 13.30 3.60a 36 4d 90 7 16.39 3.57a 38 4d 10 16.89 4.58a 38 5d VACUUM- 70 7 18.44 3.80a 22 4a,b FIR 10 19.12 4.07a 21 5a 80 7 19.95 3.55a 25 5b,c 10 18.16 4.51a 26 5c 90 7 16.72 3.19a 36 3d 10 17.81 3.63a 36 4d LPSSD 70 7 N/A N/A 80 7 21.52 2.23 27 3 90 7 24.09 1.26 28 6

N/A implies that the desired final moisture content could not be achieved. Same letters in the same column indicate that the mean values are not significantly different when considered at a confidence level of 95% (p > 0.05).

compared to chips produced only by LPSSD. This again confirms the potential benefits of the FIR-assisted drying processes in the production of highly crisp, fat-free snacks.

3.6 Concluding Remarks

The basic fundamentals of both conventional atmospheric-pressure Superheated steam drying and more novel low-pressure Superheated steam drying are presented in this chapter. The application of these processes to the drying of a wide variety of foods and biomaterials is illustrated and discussed, with emphasis placed on the use of LPSSD to preserve or even improve the properties of the drying materials.

Additional Notation Used in Chapter 3

a redness/greenness Arbitrary unit b yellowness/blueness Arbitrary unit DE overall color change Arbitrary unit L lightness Arbitrary unit S survival ratio – References 83

Subscript

0 initial

Abbreviations

CRP constant rate period d.b. dry basis FIR far infrared HAD hot air drying HPD heat pump drying IMC intermediate moisture content LPSSD low-pressure superheated steam drying SSD superheated steam drying VD vacuum drying w.b. wet basis

References

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4 Intensification of Fluidized-Bed Processes for Drying and Formulation Evangelos Tsotsas, Stefan Heinrich, Michael Jacob, Mirko Peglow, and Lothar M€orl

4.1 Introduction

Fluidized beds are widely used in industry for both the drying and wet formulation of particulate products. Wet formulation is achieved by spraying a solution or suspension (slurry) onto fluidized particles and evaporating the solvent or carrier liquid, which leads to a solid-phase formation in the form of layers (coating, granulation) or bridges (agglomeration). Consequently, formulation of the product is inherently and intimately coupled to drying in terms of both, process capacity and product structure, similar to the case of spray-drying. Due to their importance in practice, drying fluidized-bed processes have been intensively discussed in Modern Drying Technology, especially in Chapter 6 of Volume 1, where population balance equations have been presented as an instrument for the macroscopic description of wet formulation, and in Chapter 7 of Volume 3. In the latter case, the influence of material properties, operating conditions and apparatus design on the type, structure and properties of products obtained in spray fluidized-bed dryers has been discussed in detail. In the present chapter, previous discussions are extended through the aspect of process intensification. Two approaches that may significantly enhance fluidized- bed processes – namely, intensification by apparatus and flow design, and intensification by contact heating – will be treated in detail. Intensification by apparatus and flow design (Section 4.2) is based on the idea of deliberately changing the geometry and the construction of the fluidization chamber and gas entrance, which leads to so-called spouted beds. Consequently, different types of such equipment are discussed in Section 4.2.1, and details of their operating characteristics (pressure drop, operating regimes) are presented in Section 4.2.2. The intensification of gas-side mass and heat transfer that can be attained in a specific type of spouted bed equipment is quantified in Section 4.2.3. Finally, emerging possibilities for understanding the advantages or disadvantages of different types of fluidized-bed equipment by means of discrete particle modeling (DPM) are outlined in Section 4.2.4. Intensification by contact heating, which can

Modern Drying Technology Volume 5: Process Intensification, First Edition. Edited by Evangelos Tsotsas and Arun S. Mujumdar. Ó 2014 Wiley-VCH Verlag GmbH & Co. KGaA. Published 2014 by Wiley-VCH Verlag GmbH & Co. KGaA. 86 4 Intensification of Fluidized-Bed Processes for Drying and Formulation

be achieved by either immersed heating elements or by heating of the fluidization chamber wall, is discussed in Section 4.3. Following an explanation of the general principle in Section 4.3.1, the main effects and influences – as can be elucidated from experimental results – are presented in Section 4.3.2, where some modeling issues are also discussed. Since the numerous routes to process intensification presented in other chapters of this volume are also applicable to fluidized-bed drying, these discussions will not be repeated here; rather, some cross-references and short additional hints will be provided in Section 4.4, and the chapter will be closed with a brief summary in Section 4.5.

4.2 Intensification by Apparatus and Flow Design

4.2.1 Different Types of Spouted Bed

Spouted beds have a special apparatus design to create spatially correlated flow patterns of particles and gas. Whereas, gas bubbles rising through conventional fluidized beds invoke a random motion and mixing of particles, in spouted beds the apparatus design gives rise to an upwards-oriented gas jet. Velocities higher than the particle entrainment velocity force the solids to follow this jet; however, the cross-sectional area of the apparatus increases with its height in spouted beds, so that the gas velocity decreases gradually and becomes, at a certain position, too small for entrainment. Consequently, the particles fall onto relatively well-defined trajectories back to the bed, where they are recaptured by the jet and the same type of motion is then repeated. The result is that the gas velocity, which must lie between minimal fluidization and the particle entrainment velocity in conventional fluidized beds, can be much higher than the entrainment velocity in certain regions of spouted beds, but without any significant elutriation occurring from the process chamber. Due to the regularity of the imposed circulatory motion of the solids, particles will appear periodically at thesamepositionintheequipment,atwhich point either a nozzle can be placed for product formulation by coating, or a second gas inlet for rapid catalytic reaction. Moreover, the very high local values of gas velocity in spouted beds enable the stable fluidization of bulk materials that cannot be fluidized via conventional processes because of their large size and/or highly irregular particle shape. The basic design concepts of conventional fluidized beds and spouted beds are shown comparatively in Fig. 4.1. In the conventional, bubbling fluidized bed (Fig. 4.1a) the gas enters the fluidization chamber via a distributor plate at a superficial velocity that is between the minimal fluidization and particle entrain- ment velocity. Small bubbles are formed at the distributor, and these grow by coalescence while ascending through the bed. The bubble size and their rising velocity will depend on the specific two-phase system (gas and particle properties) in a manner that can be estimated according to Geldart (1973). The bubbles erupt at 4.2 Intensification by Apparatus and Flow Design 87

(a) (b)

(c) (d)

ΔPbed ΔPbed

1 2 34 1 2 3 4 56

u0 u0

Fig. 4.1 (a) Design principle and (c) bed bubbling fluidization; 4, pneumatic conveying. pressure drop of a conventional fluidized bed in Operating regions of spouted beds: 1, packed comparisonwith(b)designprincipleand(d)bed bed; 2, spout formation; 3, transition region; 4, pressure drop of a conical spouted bed. stableoperatingregion;5,pulsation;6,plugflow. Operating regions of conventional fluidized Bed pressure drop is plotted over gas velocity. beds:1,packedbed;2,transitionregion;3,stable the upper boundary of the bed, expelling particles that can be elutriated even at less than the particle entrainment velocity in the freeboard. In order to reduce elutriation, an expansion section (not shown in the figure) with an increasing cross-sectional area is usually placed on top of a conventional fluidized bed; this allows the gas velocity to be decreased such that the expelled particles can fall out from the gas flow. The ascending bubbles convey particles in their wake, thus providing an intensive mixing of solids in conventional fluidized beds. The influence of the bubbles on interface heat and mass transfer is rather negative, due to a back-mixing of the gas that reduces the driving transfer potentials and some bypass-flow action. 88 4 Intensification of Fluidized-Bed Processes for Drying and Formulation

The totally different flow pattern in a spouted bed is shown in Fig. 4.1b where, as noted above, the cross-section of the spouted bed expands with its increasing height. As a result of this geometry the gas velocity will exceed the particle entrainment velocity at the bottom, which creates the spout. However, at the top the gas velocity will be less than the entrainment velocity, so that particles can fall out and enter a regular recirculation. The typical behavior of bed pressure drop with increasing gas velocity is depicted in Fig. 4.1c and d for conventionally fluidized beds and for spouted beds, respectively. The graph for a conventional fluidized bed (Fig. 4.1c) can be subdivided into four regions, corresponding to: 1, the packed bed; 2, transition to bubbling fluidization; 3, stable bubbling fluidization; and 4, pneumatic conveying. In the case of a spouted bed, the behavior becomes more complex and depends, to a certain extent, on the type of spouted bed being considered (see discussion below). However, this dependence is usually moderate so that, according to Piskova (2002), six regions can be identified for all types of spouted bed on the graph of Fig. 4.1d. In region 1, the pressure drop increases with increasing gas flow rate, corresponding to the pressure drop of a packed bed, and reaches a maximum. The bed pressure drop is decreased again in region 2 as a spout begins to take form, creating a path of easy flow for the gas through the particles. In region 3 the spout is still very irregular, so that gas breakthrough occurs by means of both the developing spout and large gas bubbles, accompanied by strong pulsations of pressure. Beyond a certain gas velocity a stable operating regime of the spouted bed can be observed, as denoted by region 4 in Fig. 4.1d. The spout is well established and regular in this region, and the amplitude of pressure pulsations is relatively low. The average value of the bed pressure drop no longer increases with the increasing gas flow rate in region 4; rather, it remains constant at a level which is much lower than in a stably operated conventional fluidized bed of the same mass (corresponding to region 3 of Fig. 4.1c). This means that the weight of the bed, which defines the bed pressure drop in conventional fluidization, does not necessarily need to be outbalanced by the bed pressure drop of a spouted bed. Depending on the precise apparatus configuration, an excessive increase in gas velocity leads first to a new regime of pulsation (region 5) and then to a plug-like, unidirectional flow in a large part of the apparatus (region 6). The ultimate operational limit is reached when the largest particles can be entrained by the gas even in the largest cross-section of the equipment. The major goal of spouted bed apparatus design is a stable operation in region 4. In both, conventionally fluidized and spouted beds, bed porosity starts with the packed bed value in region 1 and then increases continually with increasing gas velocity in regions 2 and 3 or 2 to 6, respectively. Concerning the spatial distribution, both configurations have almost particle-free regions – the bubbles in conventionally fluidized or the spout in spouted beds – surrounded by dense regions with nearly packed bed porosity – the so-called “suspension phase” in conventionally fluidized and the region of sliding particle recirculation in spouted beds (cf. Fig. 4.1a and b). 4.2 Intensification by Apparatus and Flow Design 89

In order to fully use the major advantage of spouted beds – namely, their ability to treat very different types of particulate material – various design alternatives have been proposed and developed. An evident distinction results from the shape of the process chamber, which may be conical, conical-cylindrical or prismatic, corre- sponding to either circular or rectangular cross-sections. Solids are present only in the conical part of the equipment depicted in Fig. 4.1b which is, therefore, a conical spouted bed (Olazar et al., 1992; Olazar et al., 1993). In contrast, the solids also fill a significant part of the cylinder, set on top of the cone in the set-up shown in Fig. 4.2, so that this then belongs to the conical-cylindrical type. It should be noted that the angle and height of the cone, as well as the height of the cylindrical part, can be varied over a wide range (Mathur and Epstein, 1974). Nevertheless, the operating regions of conical and conical-cylindrical spouted beds are similar, as illustrated in Fig. 4.1d. Because of high turbulence in the spout and the regularity of particle mixing, different types of spouted bed are highly appropriate to the placement of spray nozzles for the purpose of particle formulation, such as granulation or coating. Two examples in which liquid spraying occurs directly at the entrance of the gas jet, and which results in a concurrent, upward flow of the droplets and gas with frequent droplet–particle collisions, are shown in Fig. 4.3. The equipment shown in Fig. 4.3a is of the conical-cylindrical type, whereas the cylindrical device of Fig. 4.3b is named Wurster equipment, after its inventor (Wurster, 1959). The main feature of

Fig. 4.2 Conical-cylindrical spouted bed, in contrast to the conical equipment of Fig. 4.1b. 90 4 Intensification of Fluidized-Bed Processes for Drying and Formulation

(a) (b)

Fig. 4.3 Spouted beds with circular cross-section and liquid spraying. (a) Conical-cylindrical; (b) Wurster equipment.

the Wurster equipment is a centrally placed riser tube, wherein droplets are sprayed and deposited onto the particles. Segmentation of the distributor plate with a higher permeability below the riser causes the gas flow to be significantly stronger in this tube than at the sides, thus creating a toroidal circulation of the particles. Either one-fluid or two-fluid nozzles can be used in the Wurster equipment. Unfortunately, spouted beds with a cylindrical cross-section cannot be deliber- ately enlarged, and their scale-up is difficult. However, this problem proved to be an important motivation for the development of prismatic spouted beds with a rectangular cross-section. A simple configuration of such equipment, as proposed by Mitev (1967), is depicted in Fig. 4.4a. This and other similar equipment can be simply scaled-up by prolongation, because an increase in the length of the apparatus (the direction perpendicular to the plane of Fig. 4.4) does not alter the operational regime as long as the pressure drop in the empty equipment does not

(a) (b)

Fig. 4.4 Spoutedbed with a rectangular cross-section. (a) With one gas feed (via the plenum (Mitev, 1967); (b) With two gas feeds (via the plenum and the right-hand side bottom plate (Piskova, 2002). 4.2 Intensification by Apparatus and Flow Design 91 fall below a critical value. It is not generally feasible, however, to spray in the gas jet of the equipment of Fig. 4.4a, because this jet is created on the side wall which would, as a consequence, be wetted by droplets and covered by solid deposits after drying. Rather, it is possible to construct one of the bottom plates of the fluidization chamber as an empty jacket, and to use this jacket for a second gas supply, in addition to the primary gas feed via the plenum, as proposed by Jordanova et al. (2000) and shown in Fig. 4.4b. By using this approach, rapid reactions involving the two gas streams and the particles can be conducted. In both design variants of Fig. 4.4 a porous distributor is indicated by a dashed vertical line in the gas entrance gap (enlarged post of the figure). This type of distributor is not required when the gas velocity in the gap is selected to be sufficiently high to prevent particles from falling out into the plenum. However, the ratio between the gas entrance cross-sectional area (cross-sectional area of the gap) and the surface area of the stagnant bed is an important geometric feature of both devices shown in Fig. 4.4, as noted by Mitev (1967, 1979) (cf. Eq. 4.11). Replication of the geometry of Fig. 4.4a by projection to the one or to the other side gives rise to the design configurations of Fig. 4.5. In the case of Fig. 4.5a, a particle fountain is located in the middle. Particles move to the sides at the top of the fountain, slide down the inclined bottom walls of the process chamber, and are then pushed by the gas jets back into the fountain. In this way, two symmetrically arranged rotation cells with high porosity in their horizontal axes occur over the entire length of the equipment. In the case of a gas–particle reaction (e.g., combustion) in the equipment of Fig. 4.5a, the reaction zone would be in the middle, far away from the external side walls. In contrast, in the equipment shown in Fig. 4.5b the direction of rotation of the longitudinal cells is exactly opposite to that in Fig. 4.5a, with particles rising with the two gas jets at the outer bottom plates of the fluidization chamber, moving to the middle, and then descending onto the inner bottom plates before re-entering the circulation. There is no central particle fountain in this case. The two configurations of Fig. 4.5 can be expanded for liquid spraying, as illustrated in Fig. 4.6. Spraying takes place from the bottom at the edge of the central partition in Fig. 4.6a, which means that liquid droplets are sprayed into the

(a) (b)

Fig. 4.5 Prismatic, dual-gas-entrance spouted beds which are symmetric aside their middle plane. (a)EquipmentwithcentralfountainofparticlesaccordingtoMitev(1979);(b)Equipmentwithgasjets and ascending particles at the sides according to Mitev (1979) and Piskova (2002). 92 4 Intensification of Fluidized-Bed Processes for Drying and Formulation

(a) (b)

Fig. 4.6 Expansion of the prismatic spouted beds of Fig. 4.5 for (a) bottom spray and (b) top spray of liquid.

ascending central fountain of particles. The high momentum of particles and high gas velocities in this fountain, which result in intensive drying, are rather in favor of coating or layering granulation than of agglomeration in this type of equipment. The configuration of Fig. 4.6b differs in two ways: (i) in the placement of the nozzle (it is still central, but here is employed as top-spray); and (ii) in the direction of particle cell rotation. By combining these two effects, however, liquid droplets can be sprayed onto a rather dense packing of quite slowly descending particles, and this may lead to an increased agglomerate formation. Stable rotation cells, resulting from correct equipment design, will lead to a regular, periodic passage of particles through the spray zone in both cases of Fig. 4.6. This feature may be used either to improve the uniformity of coating layers or to reduce the width of the particle size distribution of agglomerates. Many spraying (usually two- fluid) nozzles can be placed periodically along the equipment, while scale-up can be achieved by the above-mentioned prolongation and by further lateral replication, to create virtually any size. A common limitation of all types of prismatic spouted bed equipment with lengthy gas entrance gaps described so far results from the constant width of these gaps. The consequence of a constant gap width is that the pressure drop increases with the square of the gas volume flow rate, which is unfavorable in terms of automatic control. Moreover, the entrance may be blocked by solids at some places, decreasing the gas flow rate for given blower characteristics, or increasing the pressure drop and energy consumption. From these points of view, constructive solutions that would enable adjustment of the width of the gas entrance and the respective pressure drop during operation would appear desirable. Such a constructive solution has been developed by Morl€ et al. (2000) and is illustrated in Fig. 4.7. This solution is based on the use of roller valves, made from cylindrical drums which are cut flat on one side, so that rotation of the valve opens the gas entrance gap in one direction and closes the gap in the other direction, thus creating the desired adjustability of cross-sectional area and pressure drop (Gryczka, 2009). This concept has been introduced successfully to the market by the Glatt Co., under the trade name ProCell, and will be further discussed in Sections 4.2.3 and 4.2.4. It can be applied to any spouted bed with a rectangular cross-section, as illustrated by the two examples in Fig. 4.7. 4.2 Intensification by Apparatus and Flow Design 93

(a) (b)

Fig. 4.7 Prismatic spouted beds with adjustable gas inlet. (a) Equipment with central particle fountain without liquid spraying; (b) Equipment with gas jets at the sides and top spray.

4.2.2 Operating Characteristics of Spouted Beds

Several efforts to chart fluidization and the behavior of fluidized particles have been undertaken (Mathur and Ghisler, 1955; Mitev, 1967; Mitev, 1979; Li and Kwaik, 1980; Kunii and Levenspiel, 1991; Olazar et al., 1993). One such diagram, shown exemplarily in Fig. 4.8, is the modified version of a plot that was originally proposed by Grace (1986) but later expanded by information provided by additional authors and recommended for use by Piskova (2002).

Pneumatic Fast transport fluidized beds

10 ut u*

1

Spouted beds

10-1

umf Bubbling fluidized beds

10-2

A B D 10-3 1 10 d* 102

Fig. 4.8 Characteristic regions of fluidization. 94 4 Intensification of Fluidized-Bed Processes for Drying and Formulation

The modified gas velocity at the ordinate of the plot of Fig. 4.8 is defined as:

2 31 r2 3 ¼ 4 g 5 ¼ Re0 ; ð : Þ u u0 = 4 1 m r r Ar1 3 g p g g

fi where u0 is the super cial gas velocity and Re0 is the respective Reynolds number. The abscissa plots a modified particle diameter

2 31 3 r r r g 4 g p g 5 1 d ¼ d ¼ Ar3 ð4:2Þ m2 g

which is equal to the cubic route of the Archimedes number (Ar): 3 r r gd p g Ar ¼ : ð4:3Þ n2r g g Different groups of the classification of solids according to Geldart (Geldart, 1973; Molerus, 1982) can be identified on the abscissa of the plot of Fig. 4.8. However, this plot provides only a coarse estimate of the material properties and operating conditions leading to spouting fluidization, which is greatly influenced by the precise construction of the equipment and, especially, of the spout zone. In order to reliably design spouted bed equipment for practical application, the following three quantities must be known, depending on the specific configuration, for example, cylindrical or prismatic:

The minimal gas volume flow rate required for fluidization, which corresponds to the upper end of region 1 in Fig. 4.1d. The maximal pressure drop in the bed which is observed at starting fluidization, that is, simultaneously with the minimally necessary gas flow rate (highest value of pressure drop in Fig. 4.1d). The range of stable operation, corresponding to region 4 in Fig. 4.1d, especially its lower limit.

A very simple, but also coarse, method of determining the minimally required gas volume flow rate is to define it as the flow rate creating conditions of minimal fluidization at the upper surface of a stagnant bed of solids placed in the equipment under consideration. This derivation will be illustrated for the ProCell unit depicted schematically in Fig. 4.9. The superficial gas flow velocity at the free surface of the _ respective stagnant bed is the ratio of the gas volume flow rate Vg to the respective cross-sectional area, which is twofold the product of the width W0 and the length L of the apparatus, perpendicular to the plane of Fig. 4.9. By setting the superficial fl fl gas ow velocity equal to the minimal uidization velocity umf and the gas volume 4.2 Intensification by Apparatus and Flow Design 95

Vg Vg 2 2

W0

Vg Vg 2 2

Fig. 4.9 Schematic of a symmetrical ProCell unit, with two gas inlet gaps (Morl€ et al., 2000).

_ flow rate equal to its minimally required value Vg;min, the relationship _ Vg;min umf ¼ ð4:4Þ 2W0L _ is obtained. With known umf, Eq. 4.4 can be used to derive Vg;min. To this purpose, the minimal fluidization velocity can be calculated from the respective Reynolds number, which, according to Todes (Goroshko et al., 1958), is

Ar pffiffiffiffiffi Remf ¼ : ð4:5Þ 1400 þ 5:22 Ar It should be noted that the same authors derive the Reynolds number at elutriation (i.e., at the ultimate operation limit) to

Ar pffiffiffiffiffi Reelu ¼ : ð4:6Þ 18 þ 0:61 Ar In this way, the minimally required gas volume flow rate is obtained to n _ 2W0L gArpffiffiffiffiffi Vmin ¼ ÀÁ: ð4:7Þ d 1400 þ 5:22 Ar The previously explained method for the determination of the minimally required gas volume flow rate is of general applicability, but approximate. Alternatively, correlations that refer to a specific type of spouted bed may be used. A number of correlations of this type for conical and conical-cylindrical spouted beds are summarized in Tab. 4.1, based on an evaluation by Piskova (2002). The geometric quantities involved are illustrated in Fig. 4.10. All correlations of Tab. 4.1 lead to the minimal value of a Reynolds number:

ug;ind Rein ¼ ð4:8Þ ng built with the superficial gas velocity _ ug;in ¼ Vg=Ain ð4:9Þ 96 4 Intensification of Fluidized-Bed Processes for Drying and Formulation

Tab. 4.1 Correlations for the determination of the minimally required gas volume flow rate in conical and conical-cylindrical spouted beds.

Source Correlation  1:68 0:5 D0 0:57 Olazar et al. (1992) Rein;min ¼ 0:126 Ar ðÞtan g Din 0:48 1:27 2 0:57 H0 D0 Mathur and Epstein (1974) Rein;min ¼ 2:8 10 Ar Din Din 0:82 D 0:1 Romankow and Raschkovskaja (1968) Rein;min ¼ 0:364 Reelu 1 ðÞtan g ; Din Reelu from Eq. 4.6  0:1 0:25 0:59 Din H0 Nikolaev and Golubev (1964) Rein;min ¼ 0:05 Ar D0 D0  0:85 0:5 D0 1:25 Gorshtein and Mukhlenov (1964) Rein;min ¼ 0:174 Ar ðÞtan g Din

0:8 Mitev (1979) Rein;min ¼ 12:2Remf G ; 3 0:507 Remf from Eq. 4.5, G ¼ 1:09 10 Ar þ 4:1

D d

D0

γ 0 H

Din

Fig. 4.10 Characteristic dimensions of conical-cylindrical spouted beds. 4.2 Intensification by Apparatus and Flow Design 97 in the narrowest cross-section of the equipment – that is, in the gas entrance tube with the diameter Din and a cross-sectional area of Ain in Fig. 4.10. With Rein,min known from Tab. 4.1, the respective minimal velocity ug,in,min can be obtained from _ Eq. 4.8 and multiplied with Ain to evaluate Vg;min. The gas load that is at least necessary for fluidization in prismatic spouted beds can be estimated in a similar way, based on correlations from Tab. 4.2. As noted above, the second quantity necessary for the design of spouted bed equipment is the maximal pressure drop corresponding to the minimal necessary gas flow rate at the onset of fluidization, as the blower must be able to provide this pressure drop during the start-up of the process. Bed pressure drops measured by Gryczka (2009) in a ProCell apparatus are depicted in Fig. 4.11, in dependence of the gas volume flow rate. The plot shows very clearly that the maximal pressure drop in the bed is much higher – typically, two- to fourfold – than the bed pressure drop during stable operation of the spouted bed. By comparing with the more general Fig. 4.1d, it is recognized that the bed pressure drop falls very quickly after the maximal value, which means that region 2 of Fig. 4.1d is small for the ProCell unit investigated by Gryczka (2009). In contrast, the stable operating region 4 is quite larger, without any significant fluctuations of pressure drop on changing the gas flow rate. The results of investigations of the maximal bed pressure drop with different types of spouted bed are summarized in Tab. 4.3. Therein

D ¼ ðÞ e ðr r Þ ð : Þ Pf g1 mf p g H0 4 10 denotes the pressure drop that would outbalance the weight of the solids, corrected for buoyancy by the gas. Apart from the maximal pressure drop, the bed pressure drop during the stable operation of spouted beds is of interest. As demonstrated in Fig. 4.11, the latter will usually be significantly smaller. Respective literature results are presented in Tab. 4.4. The third parameter to be determined is the range of stable operation, and in particular the lower limit of this range that defines the gas volume flow rate which is at least necessary in order to reliably operate the spouted bed. A stable operation means the stability of particle circulation trajectories, which may be different depending on equipment construction (as noted above), but should not fluctuate too much in time. Investigations of the stable operating range have been conducted

Tab. 4.2 Correlations for the determination of the minimally required gas volume flow rate in prismatic spouted beds.

Source Correlation  0:6638 ¼ : H0 ðÞg 0:92 Piskova (2002) Rein;min 3 093 Reelu tan ,Reelu from Eq. 4.6 Din ¼ : 0:8 ¼ : 3 0:443 þ : Mitev (1979) Rein;min 12 2ReeluG ,Remf from Eq. 4.5, G 3 9 10 Ar 1 1 98 4 Intensification of Fluidized-Bed Processes for Drying and Formulation

Tab. 4.3 Correlations for the determination of the maximal bed pressure drop in conical, conical- cylindrical and prismatic spouted beds.

Source Equipment Correlation  2H Asenjo et al. (1977) Conical DP ¼ DP 1 þ 2:8 exp 0:156 0 max f D 0 2H0 Kmiec (1980) Conical-cylindrical DPmax ¼ DPf 1 þ 0:206 exp 0:62 D0 D ¼ ðÞ e r ; Mitev (1979) Prismatic Pmax KH0 1 mf pg K ¼ 1:2 for H0 < 180 mm 1:26H0 K ¼ 0:72 10 for H0 > 180 mm

6000

Mbed=1.98 kg

Mbed=1.45 kg 4000 M =0.95 kg a bed

P

,

d

e

b

P

Δ 2000

0 0 50 . 100 150 3 -1 Vg, m h

Fig. 4.11 Bed pressure drop as a function of gas volume flow rate for a ProCell unit according to measurements by Gryczka (2009).

Tab. 4.4 Correlations for the determination of bed pressure drop during stable operation in conical, conical-cylindrical, and prismatic spouted beds.

Source Equipment Correlation "#   D 1 D 2 Raschkowskaja Conical DP ¼ 0:45 r gH 1 þ þ ; st bed 0 D D (1969) in in : K ¼ 1 0:053 Ar"#0 13 ! 3 1 0:81ðÞ tan g 2 D0d Din Mathur and Conical- DPst ¼ Kr gH0 1 þ bed f D2 H Epstein (1974) cylindrical in 0 D ¼ ðÞ e r ; ¼ : ðÞ: þ : Mitev (1979) Prismatic Pst K1 mf bedgH0 K 0 8092 H0 0 5 D ¼ r ¼ : 0:13 Flisjuk et al. (1984) Prismatic Pst K bedgH0,K 1 0 053 Ar 4.2 Intensification by Apparatus and Flow Design 99

1000000 Onset of stable E fluidization 100000 End of stable D fluidization Ar = 2.6 107 C 10000 in B Re Ar = 1.7 105 A 4 1000 Ar = 3.2 10 Ar = 3.7 103 Ar = 1.8 103 100 1.0 10 0.1 G [%]

Fig. 4.12 Determination of the stable operating range of symmetrical ProCell equipment with two parallel gas entrances with different particles (A to E) according to Gryczka (2009). by several authors for different types of spouted bed (Mitev, 1979; Olazar et al., 1992; Piskova, 2002; Gryczka, 2009). Here, a method developed by Mitev (1979) und Piskova (2002) will be described, based on a Re-G-Ar-plot. The Reynolds number of this plot refers to the gas entrance, and is defined according to Eq. 4.8. The parameter G combines geometric features with the hold-up, defined as A G ¼ in ; ð4:11Þ Abed;0 where Ain is the cross-sectional area of the gas entrance and Abed;0 is the area of the free surface of the stagnant bed in the considered equipment. Finally the Archimedes number (Ar) expresses the material properties of the particle system (Eq. 4.3). The procedure is to vary the parameter G by variation of the hold-up and, thus, of

Abed,0 for a given piece of equipment and a certain particulate material (i.e., at Ar ¼ constant) until the limits of the stable operation range have been identified and the respective values of G and Rein have been plotted on the Re-G-Ar-diagram. The same determination is then repeated for another particulate material, that is, with another value of Ar. This procedure is illustrated in Fig. 4.12 on the basis of experimental results by Gryczka (2009) for a ProCell unit. The plot of Fig. 4.12 may be extended to include results for other types of spouted beds. In this way, Fig. 4.13 was obtained by Gryczka (2009). Reynolds numbers – as defined by Eq. 4.8 – that correspond to the lower limit of the stable operating regimes of Fig. 4.13 may be denoted by Rein,st. Correlations developed for Rein,st are summarized in Tab. 4.5. However, it should be noted that Fig. 4.13 and Tab. 4.5 provide only coarse estimates for the stable operation ranges of different types of spouted bed. The reason for this is that even small variations in the apparatus geometry – especially variations in the construction of the gas entrance – can have a significant influence on the fluid dynamics of the spouted 100 4 Intensification of Fluidized-Bed Processes for Drying and Formulation

Fig. 4.13 Working diagram for the (Piskova, 2002); E to F: Conical-cylindrical determination of the stable operation regions of spouted beds (Olazar et al., 1993); G to H: different types of spouted beds according to Prismatic spouted bed with one gas entrance Gryczka (2009). Between A and B: Conventional (Mitev, 1979); I to J: Prismatic ProCell unit with spouted beds (Mitev, 1979); C to D: Prismatic two gas entrances (Gryczka, 2009), same data as spouted bed with two parallel gas entrances in Fig. 4.12.

bed. Moreover, the presented results are mainly based on experiments conducted with near-spherical, non-sticky particles with narrow size distributions.

4.2.3 Mass and Heat Transfer in ProCell Units

In the previous sections it was shown that the flow field in spouted beds is spatially correlated, resulting in a characteristic, circulatory motion of particles; this is in contrast to the complex but, on average, essentially uniform flow patterns of conventional fluidized beds. Apart from the chamber geometry, special apparatus

Tab. 4.5 Correlations for the Reynolds number at the lower limit of the stable operating region for various types of spouted bed equipment.

Source Equipment Correlation Validity

0:8 3 0:507 3 Olazar Conical- Rein;st ¼ 12:2ReeluG 3 10 Ar < G < 1:09 10 et al. (1993) cylindrical Ar0:507 þ 4:1 0:8 3 0:443 Mitev (1979) Prismatic, one Rein;st ¼ 12:2ReeluG 0:7 10 Ar þ 1:1 < G gas inlet < 3:9 10 3Ar0:443 8 Piskova (2002) Prismatic, two Rein;st ¼ 6 10 Ar þ 1:8094 < G 1:979 7 gas inlets 42:3ReeluG < 2 10 Ar þ 6:085 0:0089 1:0486 Gryczka ProCell, two Rein;st ¼ 0:81 Ar < G < 3:7Ar 1:0486 (2009) gas inlets 190:8ReeluG 4.2 Intensification by Apparatus and Flow Design 101 design features can be used to adapt and control the flow field in spouted beds, such as air inlet openings of adjustable width in the ProCell equipment. This enables the processing of materials that can hardly be treated in conventional fluidized beds, namely large, heavy, non-spherical, polydispersed or sticky particles. However, the question arises as to whether such units can be expected to have an advantage also in terms of overall heat and mass transfer between the gas and the particles, or not. Hoffmann et al. (2011) examined this process intensification issue for a ProCell unit, and the following discussions are based mainly on these investigations. a The underlying experiment involved drying batches of water-loaded -Al2O3 fi a particles in air within the rst drying period. The -version of Al2O3 is, in contrast to the c-version of the same material, almost utterly non-hygroscopic and shows a distinct first drying period of constant drying rate. The particles used were nearly monodispersed, with a mean diameter of d ¼ 1.4 mm and a dry particle density of r ¼ 3 p,dry 1950 kg m . All experiments were conducted using laboratory-scale equipment with a cylindrical top section (diameter 150 mm; height 235 mm) and a ProCell bottom section (Fig. 4.14). The triangularly prismatic ProCell piece had a depth of 100 mm from its transparent front plane, and a height of 110 mm. The bottom angle measured 60, with the ProCell partition in the middle, and one roller valve at each side of it. As previously explained, regulation of the air inlet is performed by the

Fig. 4.14 Frontal view of the apparatus used in the measurements. From Hoffmann et al. (2011). 102 4 Intensification of Fluidized-Bed Processes for Drying and Formulation ϕ

r1

R1

h1

R h 2 r

x x0

Fig. 4.15 Geometric data of the adjustable air inlet in a scale of 3 to 1: R ¼ 15 mm; R1 ¼ 16 mm; r ¼ 10 mm; r1 ¼ 20 mm; h1 ¼ 21 mm; h2 ¼ 13 mm; x0 ¼ 3 mm. From Hoffmann et al. (2011).

two roller valves, each of which consists of a cylindrical drum that is cut flat on one side so as to create a planar surface. The vertical distance between the partition contour and the line where the plane meets the cylindrical surface of the drum is denoted by x in Fig. 4.15, and defines the width of the air entrance gap. If the drums are turned against each other, the two gaps either open or close; closing of the gap increases the local gas velocity (air inlet velocity) at a constant gas mass flow rate. The central partition unit combines the gas flows coming from the right and from the left to a spout. Symmetry of the flow field in respect to the central plane and stability of the jet are important features that must be attained by the correct design of the partition. The parameters varied in the 19 experiments reported by Hoffmann et al. (2011) fl _ were the dry solids hold-up Ms (from 68 to 291 g), the air mass ow rate Mg (from 20 to 40 kg h 1), the angular position of the gas entrance valves w (from 0 to 15; see Fig. 4.15), and the gas inlet temperature Tg,in (from 40 to 60 C). The variation of hold-up corresponds to static bed heights H0 ranging from 29 to 58 mm. The width of the gas inlet opening gap x can be calculated from the angle w to values between x ¼ 3 mm (for w ¼ 0) and x ¼ 1.38 mm (for w ¼ 15). The variation of air fl fi mass ow rate and inlet temperature leads to super cial velocities u0 in the upper, cylindrical part of the apparatus ranging from 0.27 to 0.56 m s 1. Gas velocities in 1 the adjustable inlet gaps ug,in were very much higher, ranging from 12 to 30 m s . The main measuring result of each experiment was the moisture content of the

outlet air, Yout, which remains constant over time as long as drying occurs within the first drying period. This was determined by infrared spectroscopy (as described

in Chapter 1, Volume 2 of this series). The moisture content of the inlet air, Yin, was measured in the same way. With a known temperature, Tg,in, and moisture content, Yin, of the inlet air, the adiabatic saturation temperature of the air, Tas, and the respective moisture content, Yas, can either be computed or read from the Mollier 4.2 Intensification by Apparatus and Flow Design 103 diagram. On the basis of the above-mentioned air moisture contents, the mass transfer efficiency of the equipment can be evaluated as: Y Y g ¼ out in ð4:12Þ Yas Yin This efficiency can range from g ¼ 0, which means that the air leaves the particle ¼ g ¼ system without any uptake of water vapor (Yout Yin), to a maximum of 1 for ¼ outlet gas saturation (Yout Yas). It should be noted that a distinction between wet bulb conditions at low values of g and adiabatic saturation conditions at high mass transfer efficiencies is not necessary, due to approximately equal values of wet bulb and adiabatic saturation temperatures for the evaporation of water in air. The next step of evaluation is to translate the efficiency, g, to the number of gas- side transfer units (NTU) of the considered particle system. In this context, the motion of the particles does not play a direct role, because conditions at the surface of all particles are identical everywhere in the dryer during the first drying period. However, the flow field of the gas, which is influenced by the motion of the particles, is of consequence, and it is necessary either to perform a detailed, full- scale CFD or to make assumptions concerning the gas flow. The second option has been implemented by Hoffmann et al. (2011), assuming the gas as either perfectly back-mixed or in ideal plug flow. The former corresponds to a so-called continuous stirred tank reactor (CSTR), and the latter to a plug flow reactor (PFR). In case of the CSTR assumption, the function g(NTU) is NTU g ¼ : ð4:13Þ 1 þ NTU The same dependence is described by

g ¼ 1 e NTU ð4:14Þ for the PFR limiting case. In one or in the other way, numbers of transfer units, defined as b r A ¼ g bed ; ð : Þ NTU _ 4 15 Mg fi are obtained from the measured mass transfer ef ciencies. With known Abed (total r _ fl particle surface area in the bed), g (gas density) and Mg (gas mass ow rate), values of the overall particle-to-gas mass transfer coefficient, b, and, finally, Sherwood (Sh) numbers b ¼ d ð : Þ Sh d 4 16 can be calculated (where d is the binary diffusion coefficient of water vapor in air). The results of the evaluation are presented in Fig. 4.16, where the black squares and the gray circles are Sherwood numbers calculated according to the CSTR and the PFR assumptions, respectively. The Reynolds number of the abscissa is defined fi with the super cial gas velocity in the upper cylindrical part of the apparatus, u0,to 104 4 Intensification of Fluidized-Bed Processes for Drying and Formulation

2 10 packed bed packed bed single particle single particle Sh FB Sh FB Sh app Sh app

1 (Ho min)(Ho min) 10 Sh Shapp app (Ho max)(Ho max) • PFTRPFR CSTR Sh [-] CSTR

0 10

-1 10 0 1 2 10 10 10 Re [-] 0

Fig. 4.16 SherwoodnumbersintheProCellunit,evaluatedwitheithertheback-mixing(CSTR)orthe plug flow (PFR) assumption and compared with regions for the respective Sherwood numbers of conventional fluidized beds. From Hoffmann et al. (2011).

u0 d Re0 ¼ : ð4:17Þ ng fl The variation of operating parameters other than Re0 is not re ected in Fig. 4.16. ¼ n d The Schmidt number (Sc g/ ) was almost constant for all experiments at approximately 0.66. Figure 4.16 allows for comparison of mass transfer in the ProCell unit with the mass transfer in conventional fluidized beds. To this purpose, Sherwood numbers calculated according to the well-known correlations of Gnielinski (2010) for a single particle (broken curve) and for a packed bed (solid curve) are plotted. The Sherwood numbers for conventional fluidized beds should lie in the stripe between these two lines, because the packed-bed condition corre- sponds to minimal fluidization, and the single particle condition to elutriation. In between, the Sherwood number should be almost constant, irrespective of ¼ the value of Re0,ataboutSh 9 for the particles used in the experiments (the horizontal straight line in the plot). Extensive previous analyses of interface heat and mass transfer in conventional fluidized beds (Groenewold and Tsotsas, 1997, 1999) have shown that such Sherwood numbers should be combined with the assumption of perfect back-mixing of the gas in the bed. Consequently, the value of approximately Sh ¼ 9 for the conventional fluidized bed should be 4.2 Intensification by Apparatus and Flow Design 105 compared with Sherwood numbers obtained for the ProCell unit in a similar manner, that is, by means of the CSTR assumption of Eq. 4.13. Such values are represented by the black squares in Fig. 4.16, with a range from 11.2 to 36.5 and an average of Sh ¼ 20.3, which is more than twofold larger than Sh ¼ 9. This is a clear demonstration of a very significant enhancement of particle-to- fluid mass transfer in the ProCell unit, compared to a similarly operated conventional fluidized bed. The same conclusion can be attained via the PFR evaluation. The respective data from the ProCell unit (gray cycles) are located in Fig. 4.16 significantly lower than Sherwood numbers evaluated with the CSTR assumption, because back-mixing reduces the efficiency of interface heat and mass transfer and, thus, leads to lower values of transfer coefficients, if taken into account in such coefficients instead of being considered in the balance equation of the model. For the reference case of fl “ ” the conventional uidized bed, so-called apparent Sherwood numbers (Shapp) must be used in combination with the PFR model (Groenewold and Tsotsas, 1999). By setting the right-hand sides of Eqs 4.13 and 4.14 equal to each other and expressing NTU in terms of dimensionless quantities Sh NTU ¼ AVH ð4:18Þ Re0 Sc the apparent Sherwood number can be calculated as:  Re0 Sc Sh AVH Shapp ¼ ln 1 þ ; ð4:19Þ AVH Re0 Sc where Av denotes the particle surface area per bed volume. For spherical particles it is H A H ¼ 6ð1 eÞ : ð4:20Þ V d Equation 4.20 can be applied with the expanded bed height and porosity, H and e, respectively, or with the packed bed porosity of approximately 0.40 and the stagnant bed height H0. By using the smallest and largest H0 from the experiments in ¼ Eq. 4.20 and Sh 9 in Eq. 4.19, two curves are obtained for Shapp of the fl conventionally uidized particles in Fig. 4.3. It should be noted that Shapp can be smaller than the single-particle Sherwood number. For conventional fluidized beds,

Shapp (the PFR Sherwood number) depends on H0 (it decreases with increasing H0), whereas Sh (the CSTR Sherwood number) is independent from H0 (approximately equal to 9 for the particles used by Hoffmann et al. (2011)). The difference between gray circles and the stripe of Shapp in Fig. 4.16 demonstrates the intensification of mass transfer in the ProCell unit. Based on their data, Hoffmann et al. (2011) proposed for the mass transfer in ProCell units empirical correlations of the type  p q ¼ m D D : ð : Þ Sh aRe0 4 21 H0 x 106 4 Intensification of Fluidized-Bed Processes for Drying and Formulation

Tab. 4.6 Pre-factor and exponents of the correlation for the Sherwood number (Eq. 4.21), in dependence on the assumption concerning the flow.

Assumption a m p q

CSTR 0.0468 0.57 0.48 0.79 PFR 0.0142 0.80 1.36 0.23

Although not discernible in Fig. 4.16, the parameters H0 and x were varied in the experiments. In contrast, the diameter D of the upper, cylindrical part of the apparatus was constant, and consequently dimensionless representation is the only justification for the appearance of this quantity in the correlation. The Schmidt number was also constant, and has been incorporated in the pre-factor of Eq. 4.21. The values of the pre-factor, a, and of the exponents, m, p, and q, which are listed in Tab. 4.6, depend on whether the Sherwood numbers shall be used in combination with the CSTR or with the PFR assumption in order to calculate particle-to-gas mass transfer in the unit. Nusselt numbers for gas-to-particle heat transfer can be obtained by analogy. To this purpose, the pre-factor should be split to a ¼ bScn ð4:22Þ with Sc ¼ 0.66 and an assumed value of n ¼ 1/3, then the Schmidt number should be replaced by the Prandtl number. The necessity of assuming the value of the exponent n, results from the fact that Sc was not varied in the experiments. The influence of the Reynolds number is higher in the case of plug flow (m ¼ 0.80 according to Tab. 4.6), but still existent under the CSTR assumption (m ¼ 0.57). The first finding is similar as for conventional fluidized beds, but the second constitutes a difference because Sherwood numbers of conventional fluidized beds evaluated on the basis of the CSTR assumption are nearly

independent of Re0 (the exponent m is almost zero, as discussed above). This is an indication of the significant influence of the unidirectional flow in the region of the

spout on interface transport phenomena. The static bed height H0 (i.e., the hold- up) has a great influence on Sherwood numbers evaluated by means of the PFR assumption (p ¼ 1.36 according to Tab. 4.1), but it has only small influence when evaluating for a CSTR (p ¼ 0.48). This is similar to the above-discussed behavior of conventional fluidized beds. Finally, Eq. 4.21 and Tab. 4.6 show that the mass transfer is intensified by decreasing the opening width of the valve, x, which is

equivalent to increasing the inlet gas velocity, ug,in. Hoffmann et al. (2011) attempted in their studies also a comparison between the ProCell unit and spouted beds with conventional geometry and air inlet. To this purpose, they used for conventional spouted beds on the one hand correlations by Uemaki and Kugo (1968) and El-Naas et al. (2000), and on the other hand more recent measurements and CFD simulation results (Kmiec et al., 2009; Szafran and Kmiec, 2004). The correlations indicated similar efficiencies of interface transfer in conventional spouted beds as in conventional fluidized beds, with the already discussed large advantage for ProCell units. This advantage becomes smaller when comparing with the data from Kmiec and coworkers, but it is still existent. 4.2 Intensification by Apparatus and Flow Design 107

Single-particle correlations appear to be applicable for the computation of local heat and mass transfer between the phases (Bialobrzewski et al., 2008). The resulting transfer coefficients are spatially rather uniform in conventional fluidized beds, but very unevenly distributed in spouted beds (Oliveira et al., 1998), with large maximal values in the spout. The relation between high and low heat and mass transfer regions seems to be more favorable in ProCell units than in conventional spouted beds, most likely due to extremely high air velocities in the gas entrance region of such equipment. Although the resulting process intensification must be seen in connection with pressure drop and the respective operation cost, it is an advantage by itself, as well as in terms of a smaller required apparatus volume, a smaller footprint, and less thermal exposition of heat-sensitive products.

4.2.4 Discrete Particle Modeling

Discrete particle modeling (DPM) is an advanced computational technique for particulate systems (in this case, fluidized beds) that has already been presented in Chapter 7 of Volume 3, Modern Drying Technology. DPM combines continuous (Eulerian) CFD for the gas phase with a discrete (Lagrangian) consideration of the particle phase by means of a discrete element method (DEM), and is therefore often also denoted as DEM-CFD. Its application enables the resolution of not only interactions between the gas and the particle phase, but also of particle–particle and particle–wall interactions, in the sense of a four-way coupling (compare also with Chapter 5 in Volume 1, Modern Drying Technology). The method has been recently applied by Fries (2012) for the inter-comparison of three different fluidized-bed configurations: a conventional fluidized bed with top spray (Fig. 4.17a, d); Wurster equipment (Fig. 4.17b, e); and a ProCell spouted bed unit (Fig. 4.17c, f). All three units correspond to apparatus produced by the Glatt Co. (types GF3, Wurster GF3, and ProCell 5, respectively). Some geometric data are provided in Fig. 4.17, while other data such as operating conditions, material properties and parameters of the simulations are summarized in Tab. 4.7. It should be noted that the particles were assumed to be spheres of equal size, and their number (see Tab. 4.7) corresponds to a hold- up of 0.94 kg at a particle density of 1500 kg m 3. The hold-up, fluidization gas flow rate and nozzle gas flow rate were the same for all types of investigated equipment. In order to keep the simulations tractable, droplets were not considered; however, a certain volume of each bed was defined as the spray zone. Based on results produced by Fries et al. (2011), the spray zone was assumed to be biconical, with a total length of 45 mm and a spraying angle w ¼ 40, as depicted in Fig. 4.18. The distributor plate of the Wurster equipment was segmented, with larger orifices at the center. Typical snapshots of particle positions and velocities are shown in Fig. 4.19 for the three investigated apparatuses. The top-spray fluidized bed and the Wurster equipment were cut along the central vertical plane in order to show the particles at 108 4 Intensification of Fluidized-Bed Processes for Drying and Formulation

(a) (b) (c)

Nozzle

Nozzle Air distributor (sieve plate) z y Air distributor x (perforated plate) (d) (e) (f)

250 mm 250 mm

dw = 72 mm Wurster tube 30° Nozzle Nozzle Annulus

= 200 mm Nozzle = 330 mm w h = 90 mm cone n h 180 mm 180 mm h Air inlet Air inlet Air inlet Air inlet

Fig. 4.17 Three-dimensional and cross-sectional representations of the simulated fluidized bed equipment. (a, d) Conventional fluidized bed with top spray; (b, e) Wurster equipment; (c, f) ProCell spouted bed.

Tab. 4.7 Parameters in the simulations of Fries (2012).

Parameter Value

Particle diameter 2 mm Number of particles 150 000 Fluidization air mass flow rate 360 kg h 1 Nozzle air mass flow rate 5.7 kg h 1 Gap distance below Wurster 30 mm Normal coefficient of restitution 0.8 Sliding friction coefficient of particles 0.1 Rolling friction coefficient of particles 0.05 Poisson’s ratio of particles 0.25 Shear modulus of particles 108 Pa DEM simulation time step 10 6 s CFD simulation time step 10 6 s Number of CFD grid cells 9600–17 600 Total simulation time 4 s 4.2 Intensification by Apparatus and Flow Design 109

d = 50 mm 15 mm

Llength ϕ 30 mm

Vsprayzone = 29.5 cm3 Nozzle

Fig. 4.18 Volume attributed in the simulations to the spray zone.

(a) (b)

Wurster Nozzle tube Nozzle z z y y x x

(c)

2.0

1.6

1.2

0.8

z 0.4 y x Nozzle 0.0 Particle velocity [m/s] velocity Particle

Fig. 4.19 Instantaneous particle positions and particle velocity distributions at identical process conditions.(a)Conventionalfluidizedbedwithtopspray;(b)Wursterequipment;(c)ProCellspouted bed (Fries, 2012). Colors indicate the particle velocities. 110 4 Intensification of Fluidized-Bed Processes for Drying and Formulation

the center of the fluidization chamber. The images closely resemble the particle motion patterns qualitatively discussed in Section 4.2.1, and illustrate the strong dependence of such motion patterns on apparatus design. An agitated dense bed with particle eruptions caused by bubble breakup at the top is found for the conventional top-spray equipment (Fig. 4.19a). Air from the nozzle, which was assumed to be a two-fluid nozzle, can quite strongly dilute the bed in a region located just below the atomizer, with a shape similar to that of Fig. 4.18. A dense bed with a low particle motion is seen in the lower annulus region of the Wurster equipment (Fig. 4.19b). The solids volume fraction of this region is higher than in the conventional fluidized bed. In the draft tube, particles are transported upwards and drawn towards the center of thetubebythegasflow. The gas and particle velocities are very high, due to segmentation of the distributor plate with larger orifices at the center and concurrent air injection from the atomizer. Even in the expansion chamber above the Wurster tube the gas velocity reaches 5 m s 1,comparedtoa maximum of 0.5 m s 1 in the other two devices at the same height. A relatively largenumberofparticlesarepresentinthediluteregionabovethebedinthe Wurster equipment, but no particles rise higher than 300 mm above the bottom in the other two devices. The bed is very dense at the bottom of the ProCell spouted bed (Fig. 4.19c), especially towards the inlet slots along the inclined sidewallsofthisequipment.Thespoutismuchmorediluteanddominatesthe flow pattern. There is a certain fluctuation of the spout asides the central plane of the equipment in time, so that the spout is more or less strongly tilted to one or to the other side in any snapshot picture. Selected results of the DPM study by Fries (2012) are summarized in Tab. 4.8. It can be seen that the relative velocity of collisions between particles, at 0.40 m s 1, is much higher in the spray zone of the Wurster unit (zone in front of the nozzle tip, as defined in Fig. 4.18) than in the spray zones of the two other devices (0.23 m s 1 in ProCell, 0.29 m s 1 in the conventional unit). Since high collision velocities mean that the particles will rather rebound than stick together on wet spots, it may be concluded that the Wurster equipment is more adequate for coating processes than are the other two configurations. A similar, though weaker, trend is observed when considering the relative collision velocity in the whole bed (overall value), and not only in the spray

Tab. 4.8 Selected results from the simulations of Fries (2012).

Parameter Unit Wurster ProCell Conventional

Relative collision velocity in spray zone m s 1 0.40 0.23 0.29 Relative collision velocity, overall m s 1 0.13 0.11 0.08 Particle-particle collision frequency, spray zone s 1 136 433 73 Particle-particle collision frequency, overall s 1 357 1149 1300 Collision velocity particle-wall m s 1 0.26 0.51 0.40 Time-averaged maximal energy dissipation mJ 34.9 51.5 12.6 4.2 Intensification by Apparatus and Flow Design 111

zone (relative collision velocity of 0.13 m s 1 in Wurster, 0.11 m s 1 in ProCell, and 0.08 m s 1 in the conventional fluidized bed). Certainly, coating can also be conducted by spraying from the top on a conventional fluidized bed or (and more easily) by spraying from the bottom in a spouted bed, if the operating conditions are appropriately selected for the material to be processed. However, assuming that agglomeration takes place in the spray zones of these two units, it may be concluded from the data listed in Tab. 4.8 that it would be faster in the ProCell, as the frequency of collisions between particles is much higher in the spray zone of this equipment (namely, 433 s 1)thaninthe(diluted)sprayzoneoftheconventionalfluidized bed (only 73 s 1). If applied to the entire bed, however, the same argument would be slightly in favor of the conventional unit, because of somewhat higher overall particle–particle collision frequency (1300 s 1 in the conventional, instead of 1149 s 1 in the ProCell unit). Consequently, the rate of the agglomeration process might depend on whether aggregation can take place only in the spray zones or in the entire equipment, which is a matter of drying, because the intensive drying of sessile droplets in the spray zone makes them unusable for agglomeration in the rest of the bed, whereas mild drying preserves the droplets, but would limit the equipment’s capacity. The highest particle–wall collision velocities in the ProCell (0.51 m s 1 instead of 0.26 m s 1 or 0.40 m s 1), and also highest values of maximal energy dissipated during collisions (energy dissipated by the strongest collision at every point of time, averaged over the duration of the simulation: 51.5 mJ in ProCell instead of 34.9 mJor 12.6 mJ in the two other pieces of equipment), indicate that the highest consolida- tion (lowest porosity) and the highest mechanical strength of agglomerates are produced in the spouted bed. However, this is combined with the highest risk of breakage of the as-yet unconsolidated agglomerates (which would slow the process), and with the highest risk of abrasion from coated or granulated particles. In total, the analysis by Fries (2012) indicated that spouted beds – in this case, the ProCell unit – might possess the highest versatility for the wet formulation of various products (agglomerates, coated particles, layered granules), coupled with the highest potential for intensification of the respective processes by apparatus and flow design. One important drawback of the DPM results from its high computational costs, which imposes limits on the size and number of particles that can be considered, and on the process time that can be simulated (cf. Tab. 4.7). Any expansion to consider, for example, the droplets, gas-side mass transfer (drying) or the change of particle size by the wet formulation process requires intelligent multiscale approaches. In order to obtain realistic results, parameters defining the interactions between different phases, or among the elements of one and the same phase, must be selected in an educated and elaborate manner, and the same holds for the scale- transition (coarsening) rules. However, such labor can be rewarded by insights into the influence of apparatus design on product properties and process efficiency, as noted above. 112 4 Intensification of Fluidized-Bed Processes for Drying and Formulation

4.3 Intensification by Contact Heating

4.3.1 General Principle

The heat necessary in order to evaporate the moisture from wet particles is usually provided in fluidized beds by the gas, so that this gas – typically air – acts as both the fluidizing and the drying agent. However, additional thermal energy can be inserted to the process by heating elements. The function of a heating element can be fulfilled by the jacketed apparatus wall, by some other structural part of the equipment (e.g., the partition of a ProCell spouted bed), or by panels or tubes immersed in the fluidized bed, with steam being the usual heating medium. In all of these cases, heat is transferred to the fluidized particles mainly by contact with the hot surface. This is not meant to completely replace, but rather to augment the convective heat supply from the fluidization gas, which is still present. Therefore, drying processes carried out in this way may be described as “combined convective and contact drying.” A schematic comparison of convective with combined convective and contact drying is presented in the Mollier diagram of Fig. 4.20. The benchmark for this comparison is pure convective drying along the line AC. Here, A represents the

thermodynamic condition of inlet air with a temperature of Tg,in, C is the respective condition of adiabatic saturation at a temperature of Tas, and AC is a line of constant adiabatic saturation temperature (prolongation of the fog isotherm at Tas from the oversaturated to the unsaturated region of the plot). The condition of inlet air A will change from A to C along every possible type of convective dryer, if operated adiabatically. In order to prevent condensation, and due to kinetic limitations, the condition of the outlet dryer air will usually be at a certain distance from C, which is denoted in Fig. 4.20 by point B. Projection of the vector AB on the

Fig. 4.20 Schematic comparison of convective drying (AC) with combined convective and contact drying (AE or ADE). 4.3 Intensification by Contact Heating 113 abscissa of the plot gives the change of gas moisture content in the dryer. The projection of AC expresses the maximal value of this change, that is, the specific moisture uptake capacity of the air (Yas Yin). With air of given inlet conditions as the only thermal energy carrier, the length of projection AC is unequivocally fixed. However, this projection can be prolonged by fl _ “ ” adding to the dryer a heat ow rate Qw, which is noted with the subscript w for “wall” but may be provided from any form of heating element. In a real fluidized fl _ bed, the heat ow rate Qw will be transferred directly from the heating element of temperature Tw and the surface area Aw to the bed of temperature Tbed with a fi a certain heat transfer coef cient w, obeying the relationship: _ ¼ a ðÞ : ð : Þ Qw wAw Tw Tbed 4 23 _ fl Though Qw is in reality transferred by contact to the uidized particles, in order to easily find the new trajectory of the process it might be imagined as having been fl _ fi added to the inlet air of mass ow rate Mg and speci c heat capacity cP,g, which would then increase its temperature from Tg,in to a higher value of _ ¼ þ Qw : ð : Þ Tg;in Tg;in _ 4 24 Mgcp;g The new, fictitious condition of inlet air is represented by point D in Fig. 4.20, so that the process would follow a new prolonged fog isotherm towards saturation at point F, with an exit air condition at, perhaps, point E. Alternatively, points A and E may also be connected directly with each other, whereby AE expresses the trajectory of the combined convective and contact drying process. This process is not adiabatic, so that the line AE will not be a line of constant adiabatic saturation temperature. As heat is supplied from the surround- ings – in this case from the heating elements – the line AE will be flatter than the line of constant adiabatic saturation temperature through point A, and it may even be horizontal or upwardly oriented on the plot. In any case, Tas will be higher than ð Þ Tas, and the moisture uptake capacity of the air Yas Yin will be larger than before fi (Yas Yin). Consequently, the process is intensi ed, which means that more moisture can be removed from the same amount of solids in the same dryer within less time in batch operation, or with less residence time – that is, for a higher product throughput – in continuous operation. Alternatively, smaller equipment may be used at a constant evaporation capacity and batch time or throughput. However, it might be argued that the same benefits could also be obtained by taking, in Fig. 4.20, the path ADE instead of AE. The question is then, why combine warm air with contact heat transfer, and not simply use really hot air? The answer to this point is that combined convective and contact drying can be run at temperatures Tg,in and Tw of similar and moderate magnitude, whereas the fi temperature Tg,in has to be signi cantly larger than both, Tg,in and Tw, in case of purely convective drying. Consequently, combined convective and contact drying can be carried out with a much lower value of the maximal possible temperature in the system than can pure convective drying. As the maximal possible system 114 4 Intensification of Fluidized-Bed Processes for Drying and Formulation

temperature is usually limited by specifications for products such as foods or polymers, which may suffer undesirable physical or chemical changes at too-high temperatures, drying in fluidized beds with additional heating elements has a substantial advantage for these types of thermally sensitive materials. A good example of this situation is the drying of the commodity polymer polyvinyl chloride (PVC), using extremely large-scale, cracker-close production

units. Keeping Tg,in lower than the grade-dependent threshold temperature, so as to guarantee the avoidance of thermal damage, would correspond to process AB in Fig. 4.20. As the respective water uptake capacity of the air is low, fluidized beds with a huge cross-sectional area would be required for the large-scale production of this material, which would reduce the economic competitiveness and viability of the process. An increase in the gas inlet temperature to Tg,in would be prohibited by specification, because the PVC particles or deposits would deteriorate; likewise, _ an increase in the gas flow rate Mg would not be possible because it would greatly enhance the entrainment of fines. However, the placement of steam tubes in the fluidized bed does enable processing within the necessary thermal restriction at an acceptable investment cost, which means employing a still large but reasonably sized cross-sectional area. Consequently, the processing path AE of Fig. 4.20 is viable, and this is in fact one of the two major technologies used worldwide for the drying of PVC; the second method involves the serial combination of a flash dryer with a so-called “cyclone dryer” (Heinze, 1984; Bunyawanichakul et al., 2006). Although the preceding discussion has explained, on a qualitative basis, the principles of fluidized-bed drying intensification by means of indirect heating, a detailed analysis of the process requires more precise investigation, such as the experimental and theoretical investigations conducted by Groenewold and collea- gues (Groenewold, 2004; Groenewold and Tsotsas, 2007; see also Groenewold and Tsotsas, 2001). Therefore, selected results from these authors will be detailed in the following subsections.

4.3.2 Main Effects and Influences

As the experimental investigations performed by Groenewold and coworkers were conducted at the laboratory scale, a single cylindrical immersed element was used that was heated by an electric current rather than by steam. This element was made from copper, had a diameter of 25 mm, a heated length of 50 mm, and a total length of 62 mm. In standard configuration the element was placed vertically at the axis of the (also cylindrical) fluidization chamber, which had an inner diameter of 150 mm and a height of 300 mm. The lower end of the heater was covered with a 10 mm- long poly(tetrafluoroethylene) (PTFE) cup and contacted the porous distributor plate directly; the upper end of the element was extended using a PTFE tube. In a fashion that has already been described in Chapter 1, Volume 2 of this series,

the inlet and the outlet moisture content of the air, Yin and Yout, respectively, were measured by infrared spectroscopy, so that with known mass flow rate of dry gas, _ Mg , and hold-up of dry solids, Ms, the overall water balance was used to calculate 4.3 Intensification by Contact Heating 115 both, the solids moisture content

_ ðt Mg XtðÞ¼X0 ðÞYout Yin dt ð4:25Þ Ms t¼0 and the drying rate _ r ; d _ Mg p dry m ¼ ðÞYout Yin ; ð4:26Þ Ms 6 r as functions of time. Here, X0 is the initial moisture content of the solids, p,dry is the dry particle density, and d is the particle diameter. It should be noted that the drying rate m_ (evaporation flux) refers to the total surface area of the fluidized particles. fi a Simultaneously, the heat transfer coef cient w was calculated from Eq. 4.23. To fl this purpose, the temperature in the core of the uidized bed, Tbed, was measured. Heater temperature was monitored by a Pt100 placed as close as possible to the heater surface and kept approximately constant by a PID controller. The fi temperature of the heater surface, Tw, was obtained by nite element method (FEM) computations of the temperature field in the heater, being always very close fl _ to the measured temperature. Finally, the heat ow rate Qw was determined from the electrical power consumption of the heater, after minor correction for its thermal inertia upon slight changes during temperature control. Groenewold performed these investigations in 188 experiments, all of which were in batch operation mode, and with 10 different particle fractions. However, at this point only six of these experiments will be used, performed with three particle fractions, to illustrate the influence of contact heating on fluidized bed drying. A summary of relevant particle properties and operating conditions for the data to be discussed is provided in Tabs. 4.9 and 4.10. Each of the regarded three fractions c consisted of spherical -Al2O3-particles with a narrow size distribution, and are denoted by the same abbreviation (charge name) as in the original studies (Groenewold, 2004; Groenewold and Tsotsas, 2007). As shown in Tab. 4.9, the three fractions differ from each other mainly in terms of particle size, which ranges from 1820 mm (G1800b), over 255 mm (NWA) to 49 mm (NG100). Hence, all particle sizes typically present in fluidized bed drying are covered, as are all relevant Geldart groups (D, B, and A, respectively). The operating conditions of the six selected experiments are summarized in Tab. 4.10. Of these experiments, three were conducted in pure convective mode,

Tab. 4.9 Materials in the as-discussed, selected experiments from Groenewold and coworkers.

3 Charge name d, mm rp,dry, kg m Geldart group

G1800b 1820 1070 D NWA 255 1325 B NG100 49 1850 A 116 4 Intensification of Fluidized-Bed Processes for Drying and Formulation

Tab. 4.10 Main parameters of the as-discussed, selected experiments from Groenewold and coworkers.

_ 1 1 Charge Tg,in, CTw, C Mg,kgh Yin, gkg Ms, kg

G1800b 50.2 — 65.3 7.69 0.810 G1800b 50.2 140 83.8 4.06 0.836 NWA 50.2 — 17.2 4.51 0.728 NWA 50.1 100 16.6 4.49 0.705 NG100 50.2 — 3.38 4.32 0.818 NG100 49.1 70 2.64 6.95 0.927

fi with the heater in place but switched off, so that the eld for the value of Tw is void, whilst the other three experiments represented the process of combined convective

and contact drying. The inlet gas moisture content Yin was that of ambient air, and the moisture to be removed was demineralized water. Although the gas mass flow rate and solids hold-up could not be maintained exactly the same in runs with and without contact heating, the pairs of runs compared very well one with another. This comparison was carried out in terms of drying curves (as resulting from Eqs 4.25 and 4.26), and the outcome illustrated in Fig. 4.21. The data showed that the intensification of the drying process attained by switching on the heater (empty symbols) was for every particle size very substantial with respect to the basic case of pure convective operation (full symbols), despite the relatively small ratio between the heater and bed cross-sectional area of about 2.8%. The use of a larger heating area in the bed would further increase the drying rates, or it may provide the

opportunity to decrease the temperature Tw of the heater at a constant drying rate. This was in full agreement with the qualitative discussion of the preceding section. The lower drying rates for the smaller particles in Fig. 4.21 did not necessarily mean a lower drying performance, but rather resulted from the reference of m_ to the total surface area of the solids, as noted above. With regards to the heat transfer coefficients from Eq. 4.23 for the three experiments with a switched-on heater, the data in Fig. 4.22 shows that such coefficients depend on the moisture content of the solids. With the exception of a a transient at the start of the experiments, w is increasing with the increasing solids moisture content X. This means that the heater-to-bed coefficients that would be obtained for a dry material are additionally enhanced for a wet material, which is an added benefit of the combined drying process in terms of intensification. The additional enhancement was moderate for large particles, such as those of fraction G1800b with an average particle diameter of d ¼ 1820 mm, but was stronger for smaller particles (NWA with d ¼ 255 mm) and strongest for the smallest particles. The extent can be most clearly observed in Fig. 4.23, where the relative heat transfer coefficients

awðÞX arel;expðXÞ¼ ð4:27Þ awðÞX ¼ 0 4.3 Intensification by Contact Heating 117

Fig. 4.21 Fluidized-bed drying curves of the heating.(a)Largeparticles,Gl800b;(b)Medium- three selected fractions from Tab. 4.9 for the sizedparticles,NWA;(c)Smallparticles,NG100. operating conditions of Tab. 4.10, without (full Reproduced with permission from Groenewold symbols) and with (empty symbols) contact and Tsotsas (2007). 118 4 Intensification of Fluidized-Bed Processes for Drying and Formulation

Fig. 4.22 Heater-to-bed heat transfer coefficients measured during the drying of fluidized particles ofdifferentsizeincomparisontothemodelofMartin. ReproducedwithpermissionfromGroenewold and Tsotsas (2007).

Fig. 4.23 Influence of moisture content on wall-to-bed heat transfer by means of relative heat transfer coefficients according to Eq. 4.27. Reproduced with permission from Groenewold and Tsotsas (2007). 4.3 Intensification by Contact Heating 119 defined by reference to the heat transfer coefficient of dry particles, are plotted. Here, it can be observed that heat transfer to the finest fraction NWA with d ¼ 49 mm is enhanced by more than 50% in comparison to the respective dry product. It should be noted that Macchi et al. (1999) reported an increase of heat transfer to wet particles with a diameter of 2000 mm by about 15% in a spouted bed. Brown et al. (1998) used so-called “phase-change materials” (encapsulated n-octadecane) with particle diameters ranging from 142 to 585 mm, and reported an enhancement of wall-to-bed heat transfer by up to 30% in comparison to respective inert particles. Unfortunately, another, opposing observation can also be made for the 49 mm particles (NG100) in Fig. 4.22. To this purpose, the experimentally determined a values of w were compared with values predicted using a model for contact heat transfer to fluidized particles, as devised by Martin (1984). This model (indicated by solid lines on the plot) meets the measured data at low values of particle moisture content, X, very well for the medium and large-sized particles (Gl800b and NWA, respectively), but misses them for the 49 mm particles. The main reason for this deviation is that the fine particles tend to behave cohesively and agglomerate even in the dry state, effectively behaving like larger particles, in similar fashion to particles with a diameter of 100 mm in the case of fraction NG100. It should be noted that an undesired agglomeration of fines in fluidized beds has been frequently reported in the literature (Moseley and O’Brien, 1993; see also Chapter 7 of Volume 3, Modern Drying Technology), and it has also been reported to reduce contact heat transfer (Pakowski and Mujumdar, 1982). In this respect, part of the advantage of the finest particles in heat transfer will be lost, although the remaining part may be sufficiently high for them to have the highest heat transfer coefficients (as shown in Fig. 4.22). Values of the observed heat transfer coefficients ranged from about 100 to 800 W m 2 K 1, which is typical of contact heat transfer to fluidized particles. The results of Fig. 4.22 can be further analyzed by defining a new dimensionless heat transfer coefficient by reference of the measured to the calculated heat transfer coefficient for both dry and wet particles as: a ðÞ exp X acalcðÞX ¼ 0 arel;calcðÞ¼X ð4:28Þ acalcðÞX aexpðÞX ¼ 0 and plotting it in Fig. 4.24. The horizontal basis line of this figure at arel;calc ffi 1 means that influences of particle moisture content included in the model of Martin for the contact heat transfer are the only influences present, and thus accurately captured, whereas higher values indicate additional effects which are not included in this model. Figure 4.24 shows that the former applies for the large particles (G1800b), whereas the latter is the case for the smaller particles (NWA, NG100). To better understand the physical background of the mentioned additional influences, the model of Martin should be briefly commented here, though its equations will not be given, because they are standard and well known from primary literature (Martin, 1984), handbooks (Martin, 2010), and also from Groenewold (2004) and Groenewold and Tsotsas (2007). Martin’s model is a 120 4 Intensification of Fluidized-Bed Processes for Drying and Formulation

Fig. 4.24 Transformation of the data from Fig. 4.22 with Eq. 4.28 in order to show the influence of evaporation in the particles on heat transfer. Reproduced with permission from Groenewold and Tsotsas (2007).

particle convection model, transferring ideas from the kinetic gas theory to fluidized particles that come into contact with a heating surface. When such particles are large, they remain nearly isothermal during every single contact with the heating surface, so that the only resistance to heat transfer is by conduction in the gas-filled gap between heater and the particle in short contact with it. This is the so-called “conduction-controlled” limiting case of the model. Additionally, some heat is transferred from the heating element directly to the gas, the flow velocity of which is large with large particles. At the conduction-controlled limit, it does not matter whether the particles are dry or wet. The particle moisture content will influence the wall heat transfer coefficient only via the porosity of the expanded bed, which depends on otherwise constant conditions on particle density and, thus, on the mass of water in the particles. This type of influence can be easily incorporated into the model, and it is the only significant influence of moisture content for large particles, so that the data for fraction G1800b and the predictions of the model are in excellent agreement in Fig. 4.22, lying after transformation according to Eq. 4.28 on the basis line of Fig. 4.24. In contrast, very small particles correspond to the “capacity-controlled” limit of Martin’s model. Remembering that this model has been developed for dry particles which are heated, and not for drying wet particles, it is easy to understand that particles can attain a heating surface temperature after just one wall contact when they are sufficiently small, so that the heat transfer coefficient will then result from the amount of thermal energy that can be taken up by one particle and from the frequency of particle–heater collisions. An influence of moisture content can be formally worked into Martin’s model by simply accounting for the moisture in the 4.3 Intensification by Contact Heating 121 specific heat capacity of the particles, as if the water would be heated up with the solid skeleton, but not evaporated. Along with the above-mentioned moisture- dependent particle density, this leads to the solid lines for calculated heat transfer coefficients in Fig. 4.22. As can be seen, these do not meet the measured data for the fractions NWA and NG100 very well, a point that is also illustrated by the deviation of the transformed data from the basis line in Fig. 4.24. The above-mentioned deviation is not surprising, because it would have been unrealistic to expect the mere consideration of the mass and heat capacity of the liquid to describe the data perfectly. However, it does contain an interesting and not immediately obvious aspect, which concerns the transition from the conduction- controlled region to the capacity-controlled region of the model. This transition takes place at particle sizes of about 40 mm for dry materials, so that the contact heat transfer to the fluidized bed is rather conduction-controlled above, and capacity- controlled below, this value (Martin, 1984). Drying particles with a latent heat sink in their interior certainly have a higher heat uptake capacity than dry particles, which undergo just an increase in temperature. Consequently, the influence of drying may be expected to shift the transition from the conduction-controlled to the capacity-controlled regime to smaller particle sizes (smaller than 40 mm), but not to larger particles. However, in that case the fractions NWA and NG100 with d ¼ 255 mm and d ¼ 49 mm, respectively, would be large enough for conduction control, so that they should behave in similar fashion to the very large particles of fraction G1800b, which they do not (Fig. 4.24). Hence, the conclusion is that contact heat transfer to drying particles of the above-mentioned size range is influenced by both, heat conduction through the gas-filled gap between particle and heater during a contact, and heat transfer in the interior of the particle under conditions of phase change by evaporation. The second, particle-side part of the influence makes modeling extremely demanding and tricky, because it constitutes an interrelation between contact heat transfer and drying in the sense of a higher-order, two-way coupling. Additionally, the particle model of Martin should be readjusted to longer than the originally proposed contact times between wall and particles during collisions in order to preserve the experimentally demonstrated influence of particle-side transport phenomena, even for medium-sized particles. These points will be briefly addressed in the following section, along with the more general and – from a practical viewpoint – more important task of calculating the combined convective and contact drying in fluidized beds with reasonable accuracy.

4.3.3 Further Remarks on Modeling

Groenewold, in his studies, evaluated 10 different model versions of combined convective and contact fluidized-bed drying by comparing all experimental results (Groenewold, 2004; Groenewold and Tsotsas, 2007). From this extensive evaluation, only three model versions will be briefly discussed here, denoted by the same letters as in the original studies, namely C, J, and G. 122 4 Intensification of Fluidized-Bed Processes for Drying and Formulation

Model version C is short-cut and very easy to apply. The drying kinetics of the material is described by the normalized, single-particle drying curve f(W) according to the original proposal by van Meel (1958). As already discussed in Chapter 1, Volume 2 of this series, f is the drying rate, which is made nondimensional by dividing with the constant, first period drying rate, whereas W is the moisture fi content, normalized by use of Xcr for the transition from the rst to the second drying period and Xeq for the adsorption equilibrium moisture content at the end of the process. As the normalization concept is empirical, f(W) has to be found by measurement. The drying kinetics of the single particle is then scaled-up to the fluidized bed by an appropriate model. In conjunction with the function f(W), this model does not need to consider explicitly the energy balances, so that the simple formulation introduced by Groenewold and Tsotsas (1997) is sufficient. In this, the particles are ideally back-mixed, bubbles are considered to rise through the bed in plug flow, exchanging heat and mass with the suspension phase (dense phase). Plug flow is also assumed for the gas flowing through the suspension phase, which is primarily responsible for the drying, but this assumption is corrected by incorporating the influence of suspension gas back-mixing into apparent Sherwood numbers for particle-to-gas mass transfer. The previously stated Eq. 4.19 describes such apparent Sherwood numbers. All model parameters are taken from the literature, so that there is no degree of freedom for fitting. Contact heat transfer is accounted for in model version C by artificially increasing the temperature of the fl inlet air to Tg,in according to Eq. 4.24. To obtain the necessary heat ow rate from fi a Eq. 4.23, the wall heat transfer coef cient w is calculated by using the original model of Martin. Herein, the influence of moisture content on particle density and heat capacity is taken into account, while the previously discussed, “second-order” influences of phase transition in the particles during heat transfer are simply fi a ignored. Consequently, the heat transfer coef cient w has exactly the values shown by the solid lines in Fig. 4.22. Drying curves computed in this way correspond to the solid lines in Fig. 4.25. As Fig. 4.25a shows, the agreement with the measured data is excellent for fraction G1800b. However, for fraction NWA more deviation occurs between measurement and prediction (Fig. 4.25b), because the contact heat transfer coefficients cannot be described exactly by Martin’s model for these particles (Fig. 4.22), due to the neglected influence of the latent, evaporative heat sink. Especially, the slope of the measured drying rates with decreasing moisture content is not captured exactly; rather, it is underestimated at high values of X, and overestimated at low values. However, the prediction result is still usable from a practical point of view, which means that the intensification caused by contact heat transfer over the entire duration of the process is evaluated quite accurately. Because of the identified effect of agglomeration, no comparison will be made for fraction NG100. The results of model version J are also depicted (as broken lines) in Fig. 4.25. This model version is similarly simple and almost identical to model version C, apart from the fact the heat transfer coefficient from the heating element is now calculated by assuming an infinitely large specific heat capacity of the particles in Martin’s model. This is equivalent to saying that the evaporative sink in the 4.3 Intensification by Contact Heating 123

Fig. 4.25 Prediction of drying curves by the short-cut model versions C and J for (a) large particles, G1800b, and (b) medium-sized particles, NWA. Reproduced with permission from Groenewold and Tsotsas (2007). particles is so strong as to prevent any penetration of heat into them, irrespective of the particle size, and removing completely the capacitive limitation from the model of Martin. For coarse particles, the result of model version J is the same as the result of model version C, and can thus not be distinguished graphically in Fig. 4.6a, because heat transfer to such particles is anyway controlled by conduction 124 4 Intensification of Fluidized-Bed Processes for Drying and Formulation

Fig. 4.26 Prediction of drying curves by sophisticated model versions, especially model version G, for medium-sized particles, NWA. Reproduced with permission from Groenewold and Tsotsas (2007).

in the gas-filled gap to the heater, and not by the heat uptake capacity. For the medium-sized particles (Fig. 4.6b), the drying rates predicted by model version J are somewhat higher (though not by much), because even such particles are located rather in the conduction-controlled region according to the original version of the model, as previously discussed, so that the removal of the capacity limitation does not play a crucial role. The best prediction of the drying curve of fraction NWA was achieved by Groenewold and colleagues with model version G. The results of this model version correspond to the upper, somewhat shaky curve in Fig. 4.26, and they are in excellent agreement with the experimental results for particle moisture contents larger than about X ¼ 0.04. The price for the better agreement is, though, a much more complex structure of model version G. In this version, single-particle drying is modeled, rather than by normalization, by a numerical solution of the volume- averaged mass and heat transfer equations, similar to the equations presented in Chapter 1, Volume 1 of this series (see also Whitaker, 1977). Numerical instabilities created the above-mentioned fluctuations of the solution. With regards to the fluidized particle system, an explicit consideration of the energy balances is now necessary, and this corresponds to an upgrade of the approach of Groenewold and Tsotsas (1997) published by Burgschweiger and Tsotsas (2002), which is more sophisticated and profound than the original, but also more difficult to implement. Moreover, in order to be able to harvest the effect of the coupled consideration of heat transfer and drying in the heat transfer coefficient, the model of Martin must be modified to larger particle–wall contact times during collisions (cf. Molerus et al., 1995). In doing so, the good agreement of this model with a large amount of 4.3 Intensification by Contact Heating 125 heat transfer data for dry materials should not be lost, and a compensatory correction in the gas–gap conduction term is therefore necessary. However, the extent of this compensation has certain limits (Groenewold and Tsotsas, 2007), and its physical justification is not evident, so that the heat transfer part of the model might be considered as overmanipulated. In combination with the high computa- tional demands of model G, it may be concluded that the short-cut versions C and J are more appropriate for practical tasks. It should be noted that the here undiscussed model versions, such as A and D from Fig. 4.26, include some interesting elements. In some cases, the modified normalization of single-particle drying kinetics as proposed by Burgschweiger et al. (1999) was used. This enabled a separation of the influences of water vapor pressure decrease by adsorption and intraparticle transport resistances – which are intermingled in the van Meel approach – from each other. It requires use of the full fluidized-bed drying model of Burgschweiger and Tsotsas (2002), but makes sense c for highly hygroscopic materials, such as -Al2O3. In another version, the particle model of Martin was replaced with the packet model of Kunii and Levenspiel (1991) fi a for calculating the contact heat transfer coef cient w. The latter is based on the penetration of heat in assemblies of many particles (“packets”) in contact with the heating surface for a certain time, being similar to contact drying models (Tsotsas et al., 2007; Kwapinska et al., 2008). However, none of the additionally investigated versions was significantly more successful in predicting the measured data than those discussed, nor had any grave advantages from a theoretical point of view. With regards to the many additional experiments conducted by Groenewold, a series performed with a horizontal arrangement of the heater should be mentioned. In these runs, the same heater was located in the middle of a diagonal of the bed, with a distance of 70 mm between its axis and the distributor plate. Both heater ends were insulated using PTFE cups. For particles larger than 770 mm, the measured contact heat transfer coefficients and drying curves were approximately the same as in case of the vertically placed heater, whereas somewhat smaller heat transfer coefficients were obtained with the horizontal arrangement for particles smaller than 255 mm. It should be noted that none of the 10 investigated particle fractions had an average diameter in the range of 255–770 mm. Results on the influence of the arrangement of horizontal tubes on bubbling and heat transfer in large fluidized-bed dryers have been reported by Jonassen (1999), though for a limited selection of particulate materials (potato cubes were the only wet material used). In total, the main results of the comprehensive investigations performed by Groenewold can be summarized as follows:

A very significant intensification of fluidized-bed drying can be attained by using immersed heating elements or heated walls. The drying process can be computed quite easily and accurately for particles larger than 770 mm. A coupling between contact heat transfer and particle drying occurs for particles smaller than 255 mm. This is additionally beneficial in terms of process 126 4 Intensification of Fluidized-Bed Processes for Drying and Formulation

intensification, but difficult to simulate. Relatively simple models still provide acceptable results. Complex models are somewhat more accurate, but cumber- some in their implementation and in need of further development in their theoretical details, perhaps in combination with CFD (cf. Kuipers et al., 1992; Gustavson and Almstedt, 2000; Wang and Rhodes, 2003). Very small particles of about 50 mm may tend to agglomerate, even in a dry state, which would reduce the beneficial influence of contact heating.

4.4 Further Methods of Intensification

The principle of fluidized-bed drying intensification, as described in Section 4.3, was to combine the energy supply that takes place by means of the fluidization gas with additional energy supplied by contact of the particles with heated surfaces. In a similar way, other methods of supplementary energy input can also be used, such as microwave heating or heating by infrared radiation. Water in the material to be dried can be liquid, or it can be frozen, as in fluidized-bed atmospheric freeze- drying. Although a variety of hybrid processes have emerged in this way they will not be discussed at this point; rather, they are described in detail in other chapters of this volume of Modern Drying Technology (notably Chapters 5, 9, and 10). It should be noted that, irrespective of the energy source used, this energy can be directly provided to the fluidized particles to be dried, or it can be supplied first to inert, typically large particles, which transfer then the heat in their turn to the (typically small) product particles. In such a process the large particles take on the role of the immersed heating elements of Section 4.3, the difference being that they are not fixed in space but are mobile. Their own source of energy can derive from the hot fluidization gas (i.e., convective), from hot walls or inlets (by contact), via dissipation of microwaves, or from a radiant emitter. As already mentioned, these pathways are treated either in the present chapter or in chapters elsewhere in this volume; however, one possible and interesting track that is not discussed elsewhere is briefly introduced here, namely heating by induction. The pathway of induction means that electric energy is applied to create a magnetic field in the fluidization chamber; the magnetic field, in turn, creates an electric current in large and inert metallic particles, which are co-fluidized with the product. The energy of this electric current is immediately dissipated as heat. Such heat is transferred to the fluidization gas and then from the gas to the particles of the product to be dried; alternatively, it may be transmitted directly to the product particles by contact (by collisions between inert and product particles). In order for this process to function, the inert metallic particles and wet product particles must fluidize in a rather homogeneous fashion – that is, their fluidization properties must match. The Archimedes number (see Eq. 4.3) shows that fluidization behavior depends on both the diameter and density of particles; consequently, compact metallic particles will never have the same fluidization behavior as would, for example, food particles, unless they are much smaller. However, as it is 4.5 Conclusion 127 preferable to use larger rather than smaller inert particles, they must have a lower density than the product, but be metallic in nature. This is achieved by using hollow metallic spheres that can be fabricated with different wall strengths so as to enable the above-mentioned matching. The first investigations into induction (Stresing et al., 2011, 2013) showed the process to be extremely efficient in terms of heat transfer, with very high gas temperatures being attained very soon after switching on the induction field. A fast thermal response occurring after an essentially instantaneous change of electrical input can provide new opportunities for the automatic control of drying processes. It has also been shown that the magnetic field can be easily and successfully focused on the fluidization chamber, with very small losses to the environment. Clearly, continuing studies in this area are expected to establish a new and innovative alternative for the enhancement of fluidized-bed drying by induction heating.

4.5 Conclusion

With heat transfer from gas to particles occurring in all fluidized-bed drying processes, additional heat can be supplied in a variety of ways, either directly to the product or via cofluidized inert particles. Moreover, both the fluidization chamber geometry and air inlet design can be varied in many different ways, for example replacing conventional with spouted fluidization. The hybridization of energy supplies and apparatus design can contribute to significant process intensification, which in turn can lead to higher production capacities using constant sizes of equipment and thermal exposition, or to smaller equipment with less thermal exposure of the product at a constant throughput. Combination of these factors can provide rich opportunities for improved and new industrial processes of drying and wet formulation, and also present manifold challenges and opportunities for research and scientific development.

Additional Notation Used in Chapter 4

G ratio of air inlet to bed surface area – x width of air entrance gap m

Greek Letters c bottom half-angle of cone rad w angular position of gas entrance valve rad w half-angle of spray cone rad 128 4 Intensification of Fluidized-Bed Processes for Drying and Formulation

Subscripts and Superscripts

app apparent elu elutriation f corresponding to bed weight st stable operation 0 initial, superficial, reference, static (packed bed) fictitious gas inlet condition, modified variable

Abbreviations

CFD computational fluid mechanics CSTR continuous stirred tank reactor DPM discrete particle modeling FEM finite element method NTU number of transfer units PFR plug flow reactor PID proportional integral derivative PTFE poly(tetrafluoroethylene) PVC poly-vinyl-chloride

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5 Intensification of Freeze-Drying for the Pharmaceutical and Food Industries Roberto Pisano, Davide Fissore, and Antonello A. Barresi

5.1 Introduction

Commonly, process intensification entails the development of new technologies (i.e., methods and/or equipment) that, compared to those traditionally used today, show the way towards a dramatic improvement in manufacturing. Process intensification is therefore a strategy used by companies to enhance production in terms of reductions in equipment size, increases in the production capacity within a given equipment volume, decreases in energy consumption, reductions in wasted energy and also in the costs associated with waste products. The final result of process intensification is therefore the development of smaller, cleaner, and energy- efficient technologies (Stankiewicz and Moulijn, 2000). When freeze-drying pharmaceutics and foods, industry strives to develop new technologies in order to reduce drying times and hence energy consumption. Typically, these new technologies are product-specific, as requirements for process intensification may vary with the original physical state of products being freeze- dried. In this regard, products can be classified into three categories: liquids; individually quick-frozen products; and a combination of these. In general, in the pharmaceutical industry the materials being dried are solutions, whereas the food industry treats both liquids and solids. Process constraints may be different for the pharmaceutical and food industries, as well as for the above types of product. It follows that a process modification can be beneficial or applicable only to specific material types or equipment. For example, in the pharmaceutical industry sterility is often a predominant require- ment, while minor attention is given to energy efficiency. Therefore, a new technology that enhances heat and mass transfer, but compromises sterility conditions, is not suitable for the drying of drugs. Both, the freezing and the drying processing steps must be considered for process intensification (Elia and Barresi, 1998). For liquid products, new methods have been introduced to control freezing, as this impacts on product characteristics and, subsequently, the drying phase. In addition, automatic control systems were

Modern Drying Technology Volume 5: Process Intensification, First Edition. Edited by Evangelos Tsotsas and Arun S. Mujumdar. Ó 2014 Wiley-VCH Verlag GmbH & Co. KGaA. Published 2014 by Wiley-VCH Verlag GmbH & Co. KGaA. 132 5 Intensification of Freeze-Drying for the Pharmaceutical and Food Industries

designed for the drying stage (Pisano et al., 2010; Barresi and Fissore, 2011) in order to maximize the rate of water removal, and to reduce drying times and, in turn, the processing costs. It must also be noted that freezing has its own time demands, which can also be reduced. In this respect, unconventional methods have been investigated in order to intensify the freezing of both solids and liquids (Elia and Barresi, 1998). Evaporative freezing is one of these methods, and when Ghio et al. (2000) compared the evaporative and conventional freezing of vegetables they showed that pressure reduction would allow a rapid freezing of the material being dried, as well as a partial evaporation of water (ca. 20–40% of the total). This phenomenon makes it possible to reduce both the time and the energy consumption for the subsequent drying step. However, the use of this method is limited as it may damage the product due to cracks being formed during the depressurization. While the pharmaceutical industry produces high-value products which can justify the high costs of freeze-drying, the food industry produces lower-value products and has to face large production volumes. For this reason, new methods have been proposed for food manufacturing as a valid alternative to vacuum freeze- drying. Such an example is the atmospheric freeze-drying that, as shown in this chapter, allows a dramatic reduction in energy consumption. Of course, any new method implies the development of novel and, perhaps, unconventional equip- ment, and atmospheric freeze-drying has not escaped this trend. New techniques using alternative energy sources, such as ultrasound and microwave, were employed to further accelerate the rate of drying and improve the energy efficiency of the process. For this purpose, hybrid operations were also proposed that involve more drying techniques and can lead to high-quality product and faster drying (Venkatachalapathy and Raghavan, 1999). Although process intensification mainly focuses on novel methods and equipment, it can also entail the rediscovery of established technologies, which were applied on the laboratory scale or developed for a different application. A typical example is the continuous freeze-drying plant (Rey, 2010), a technology that has been widely used in food industry since the 1960s and has led to larger productivity and more homogeneous products. In the biological and pharmaceu- tical industries, freeze-drying is still a batch operation and, despite the elaborate equipment design and sophisticated control systems recently introduced, the heterogeneous drying behavior within a batch – as well as from batch to batch – remains a problem of great concern. One potential solution to this problem is to shift from batch to semi-batch or continuous operation, as did the food industry many years ago. In this chapter, attention is focused on the challenge of process intensification in freeze-drying. The problem of process optimization is first addressed, and exergy analysis is used to identify the best operating conditions to be implemented during drying. Some of the methods and technologies proposed to date to reduce drying times and leading to a more sustainable process are then described and discussed. The advantages and disadvantages of the above methods, as well as their practicality for different types of material, are also outlined. 5.2 Exergetic Analysis (and Optimization) of the Freeze-Drying Process 133

5.2 Exergetic Analysis (and Optimization) of the Freeze-Drying Process

The amount of energy required by a freeze-drying process is very large – in fact, it is almost double that used in a conventional drying process (Flink, 1977). The most critical step is the primary drying, as this accounts for about 45% of the total energy required by the process (Ratti, 2001). In order to save energy and maximize productivity, besides preserving product quality, one possible strategy consists of minimizing the duration of the process: this requires an identification of the values of temperature of the heating shelf

(Tshelf) and of pressure in the drying chamber (Pc) that make it possible to minimize the drying time, besides satisfying the constraints on the maximum temperature allowed by the product, and on the maximum sublimation flux allowed by the freeze-dryer. This problem can be solved either in-line or off-line (Pisano et al., 2013). In the first case, it is necessary to use a monitoring and control system (Barresi et al., 2010; Pisano et al., 2010; Barresi and Fissore, 2011): mathematical modeling is required to estimate the state of the system on the basis of the available measurements (Barresi et al., 2009; Fissore et al., 2011b) and to calculate the values of the manipulated variables (with the goal of achieving the aforemen- tioned goals). In the second case, mathematical modeling can be used to calculate off-line the design space of the primary drying, thus determining the optimal values of the operating conditions (Giordano et al., 2011; Fissore et al., 2011a). The exergy analysis of the process can be a powerful tool to identify the best operating conditions. Exergy is defined as the maximum amount of work that can be extracted from a physical system by exchanging matter and energy with large reservoirs at reference states. The exergy is thus a measure of energy quality, and can be associated with the irreversibility occurring during the process. In order to optimize the freeze-drying process from the energetic point of view, it is thus required to minimize the exergy losses; this would improve the economic efficiency (and the sustainability) of freeze-drying by increasing the efficiency of energy utilization in the process. Few reports have been made in the scientific literature regarding the exergy analysis of freeze-drying processes. Liapis and Bruttini (2008) provided a detailed exergy analysis of the various steps of a freeze- drying process (freezing, primary and secondary drying), as well as of vacuum pumping and of vapor condensation, by using a detailed two-dimensional model, while Liu et al. (2008) used a one-dimensional model to carry out a similar analysis. In order to combine information obtained through the design space of the process with the exergy analysis, a simplified approach is required: many calculations are necessary, as the exergy analysis needs to be repeated for all the potential values of Tshelf and Pc that could be used to carry out the process, so that computational time becomes an important issue. Fissore et al. (2014) used the simplified one-dimensional model of Velardi and Barresi (2008) to calculate the temperature profiles in the frozen layer and the sublimation flux as a function of the operating conditions, while the temperature profile in the dried layer is calculated solving the Fourier equation written for the dried cake. The approach of 134 5 Intensification of Freeze-Drying for the Pharmaceutical and Food Industries

mass flux drying x = 0 chamber

dried product

x = Ld interface of sublimation frozen product

x = L0 shelf heat flux

Fig. 5.1 Sketch of a freeze-drying process during the primary drying stage.

Liu et al. (2008) can be used to calculate the exergy losses, but some assumptions of these authors have been removed as their validity appears to be questionable; for example, that the vapor pressure at the sublimation interface is equal to the chamber pressure. The exergy losses due to heat transfer during primary drying pd fi (Exloss;h) are calculated from the temperature pro le in the dried layer (see Fig. 5.1): ð ð td Ld l 2 pd ¼ T0 d dT ; ð : Þ Exloss;h A 2 dxdt 5 1 0 0 T dx

where A is the area of the sublimation interface, T0 is a reference temperature, T l the temperature of the product at axial position x, d and Ld are respectively the thermal conductivity and the thickness of the dried cake, and td is the duration of primary drying. The exergy input resulting from heat transfer above the reference temperature pd (Exin;h) is given by: ð td pd dT T 0 Ex ; ¼ A ld 1 dt: ð5:2Þ in h ¼ 0 dx x¼0 T x L0

Assuming that the reference temperature T0 is well below product temperature during primary drying, then the exergy input due to heat transfer is given by: pd ¼ Ti Ti Tfreeze þ D T0 þ Exin;h T0mwetcp;f ln mwet hsub 1 Tfreeze T 0 Ti T 0 T 0 T i þ T 0 mw;in mw;fin cp;d ln ; T i T 0 ð5:3Þ

where mwet is the mass of product, mw,in and mw,fin are respectively the mass of water in the product at the beginning and at the end of primary drying, Tfreeze is the product temperature at the end of the freezing stage, and Ti is the product temperature at the sublimation interface. With respect to mass transfer in the product being dried, the presence of noncondensable gases resulting from leakage and outgassing of the surfaces in the 5.2 Exergetic Analysis (and Optimization) of the Freeze-Drying Process 135 system, as well as inert gases from the material being dried, can be neglected. fl pd Thus, the exergy losses due to water vapor ux in the dried layer (Exloss;m) are given by: ð ð ~ td Ld ðÞ pd ¼AT 0R 1 dpw x ; ð : Þ Exloss;m ~ jw ðÞ dxdt 5 4 Mw 0 0 pw x dx

fl where jw is the sublimation ux, and pw is water vapor partial pressure at axial pd position x. The exergy input in the dried layer due to mass transfer (Exin;m) is given by: ð td ~ p pd ¼ T i T i þ R w;i : ð : Þ Exin;m AT 0 jw cp;v 1 ln ~ ln dt 5 5 0 T 0 T 0 Mw P0 The vapor produced in the drying chamber arrives onto the cold surface of the condenser. The change in temperature, as well as in the state, of the vapor in the condenser can be described as follows:

the temperature of the vapor (Tin,cond) is cooled down to the temperature that causes desublimation of vapor (Tdes); the vapor desublimates at the temperature Tdes; fi the temperature of the ice decreases to reach the nal value (Tout,cond).

The energy input of the vapor condenser is given by: ð td ¼ þ D þ : ð : Þ Qcond A jw Tin;cond T des cp;v hs Tdes T out;cond cp;ice dt 5 6 0 cond Then, the exergy input (Exin ) can be calculated as: cond ¼ T0 ; ð : Þ Exin Qcond 1 5 7 T cooling

cond where Tcooling is the temperature of the cooling medium. The exergy losses (Exloss ) can be calculated using the following equation: cond ¼ T des T des Tcond;in þ Exloss mw;in mw;fin T 0cp;v ln T cond;in T cooling 1 1 Tout;cond T out;cond T des þ T0Dhs þ T0cp;ice ln T cooling T des T des Tcooling : ð5:8Þ The vacuum pump is used to reduce pressure in the drying chamber during the start-up stage of the process, and to evacuate the noncondensable gases resulting from leakage (air), from controlled leakage used for pressure control (nitrogen, in most cases), and from the material being dried. The exergy input of the vacuum system is given by the power input multiplied by the drying time. In order to calculate the power input it is necessary to know the flow rate of gases that are 136 5 Intensification of Freeze-Drying for the Pharmaceutical and Food Industries

(a) (b) 70 70 60 60 50 50

, % 40 40

loss 30 30 Ex 20 20 10 10 0 0 5 10 15 20 5 10 15 20 chamber pressure, Pa chamber pressure, Pa

Fig. 5.2 Values of the exergy losses (given as percentage of the total exergy losses) for different values of chamber pressure in the primary drying stage in the drying chamber (&), for vapor condensing ( ) and vacuum pumping (&). (a) Tshelf ¼20 C, (b) Tshelf ¼ 0 C.

pumped out. The exergy losses due to the compression of a perfect gas can be written as: ~ pump ¼ R Pout ; ð : Þ Exloss mg ~ T 0 ln 5 9 Mg Pcond

where mg is the mass of noncondensable gases evacuated by the pump, given by the product of the gas flow rate due to leakage (that can be determined by means of ad hoc tests for freeze-dryer qualification) and to controlled leakage (that can be measured by means of a mass-flow meter) and drying time. Figure 5.2 shows the values of the exergy losses in the primary drying stage considering the contribution in the drying chamber for vapor condensing and vacuum pumping (given as percentage of the total exergy loss), for different values of chamber pressure. Here, the freeze-drying of coffee extract is considered as this product is the most common freeze-dried liquid in food industry, due to the fact that the process preserves the coffee flavor to a great extent. It emerges that, when the value of the shelf temperature is low, then most of the exergy losses occur in the condenser and the contribution of the other two sources depends on the chamber pressure. However, when the value of the shelf temperature is increased, the contribution of primary drying in the chamber becomes much more important, especially at lower values of chamber pressure, while the exergy losses of the vacuum pump become almost negligible. Figure 5.3 shows the absolute values of the principal exergy losses in the primary drying step as a function of chamber pressure and heating shelf temperature. It appears that, while the shelf temperature seems to have no effect on the exergy losses in the condenser (this is due to the different contributions of the phenomena occurring in the condenser on the exergy losses), its effect is remarkable on the drying chamber losses. In particular, the higher the temperature of the shelf, the higher are the exergy losses. This is due to the fact that a higher shelf temperature causes the temperature gradient in the dried layer to be larger. 5.2 Exergetic Analysis (and Optimization) of the Freeze-Drying Process 137

(b) (a) 200 350

150 300 , kJ/kg , kJ/kg , 100 loss loss Ex Ex 250 50

0 200 5 10 15 20 5 10 15 20 chamber pressure, Pa chamber pressure, Pa

Fig. 5.3 Values of the exergy loss for the primary drying step in (a) the drying chamber and (b) for vapor condensing as a function of chamber pressure and heating shelf temperature (solid line: 30 C; symbols: 0 C).

The chamber pressure affects both exergy losses in the drying chamber and in the condenser, but the effects are different: while an increasing chamber pressure decreases the exergy losses in the chamber, it increases the exergy losses in the condenser. The chamber pressure affects both heat transfer and mass transfer in the dried layer of the material: the temperature gradient can decrease due to the heat transfer enhancement, and this may reduce the exergy losses. On the other hand, when the chamber pressure increases, the pressure gradient – and thus the exergy losses due to mass transfer – are increased. In a heat transfer-controlled process, heat transfer in the product plays a more important role in the exergy losses than mass transfer in the material does, and so the exergy losses decrease when the chamber pressure is increased. In a mass transfer-controlled process, the effect of chamber pressure would be the opposite. As the process investigated is under heat transfer control, exergy losses decrease when increasing the chamber pressure. With respect to the condenser, when the chamber pressure is increased so too is the sublimation temperature, and this is responsible for the increase in exergy losses. The effect of the operating conditions on the cumulative exergy losses of the process is shown in Fig. 5.4. It appears that the exergy losses can be minimized by working at a high shelf temperature and a low chamber pressure. Previous results should be considered when the design space method is used to develop a freeze-drying cycle. Figure 5.5 shows the design space for coffee extract processed on trays in a pilot-scale freeze-dryer. The approach of Fissore et al. (2011a) has been used to calculate the design space. It appears that the design space is affected by the dried layer thickness as it changes the resistance to vapor flow.

Consequently, a couple of values of Tshelf and Pc that belongs to the design space at the beginning of primary drying may lie outside the design space when drying continues. If a simple cycle is desired, where the values of the operating conditions are not modified during the primary drying, then it is necessary to consider the 138 5 Intensification of Freeze-Drying for the Pharmaceutical and Food Industries

(a) (b) 800 800

700 700

600 600 , kJ/kg , kJ/kg loss loss 500 500 Ex Ex 400 400

300 300 -30 -20 -10 0 5 10 15 20 shelf temperature, °C chamber pressure, Pa

Fig. 5.4 Values of the total exergy loss during the primary drying step as a function of (a) shelf temperature (solid line: 5 Pa; symbols: 20 Pa) and (b) chamber pressure (solid line: 30 C; symbols: 0 C).

curve of the design space calculated for a value of dried layer thickness approaching the total product thickness (Fig. 5.5a). Obviously, there are many values of operating conditions that make it possible to fulfill a certain constraint about the maximum product temperature. In order to choose among these, it might be worthwhile considering the values of the sublimation flux, with the goal of reducing the duration of primary drying. Figure 5.5b shows values of the sublimation flux in a certain time instant during primary drying (the trend of these curves is the same throughout the primary

(a) (b) 0 0.6 -5 0.5 L /L 0.01 d 0 -10 0.5 0.4 -15 0.99

j , kg h-1m-2 w

shelf temperature, °C -20 0.25

-25 8642 201816141210 8642 201816141210 chamber pressure, Pa chamber pressure, Pa

Fig. 5.5 (a)Designspacefortheprimarydrying L0 ¼ 0.5 (dashed line identifies the design stage of a coffee extract freeze-drying process space).Theproductisloadedontheshelfusinga calculated for different values of dried cake metallic tray; the thickness of the frozen product thickness; (b) Values of the sublimation flux as a is equal to 12.5 mm and the solute mass fraction function of the operating conditions when Ld/ is 25%. 5.3 Process Intensification in Vacuum Freeze-Drying of Liquids 139

Fig. 5.6 Valuesoftheexergyyieldasafunctionoftheoperatingconditions(dottedcurves).Solidline identifies the design space when Ld/L0 ¼ 0.5. drying stage): the optimal operating conditions (i.e., those that maximize the fl sublimation ux) correspond to low values of Pc (e.g., 8 Pa in this case) and high values of Tshelf (0 C in this case). This information about exergy losses can be used to calculate an exergy yield (dividing the exergy losses by the exergy input), and these data can be added to the design space of the product. An example of these calculations is shown in Fig. 5.6 for the case study of the coffee extract freeze-drying process. It emerges that the values of Tshelf and Pc that allow the maximization of exergy yield are, in this case, those corresponding to high values of chamber pressure and low values of shelf temperature. Hence, when designing the freeze-drying cycle a compromise must be achieved, as the goals are to minimize the duration of primary drying (and, thus, to maximize the sublimation flux) and to maximize the exergy yield (i.e., to minimize the exergy losses).

5.3 Process Intensification in Vacuum Freeze-Drying of Liquids

In the previous section, exergy analysis was applied to the freeze-drying process in order to identify the optimal values for processing conditions to be used during primary drying. Here, the problem of process intensification is addressed from a different point of view – that is, to reduce the cycle time by modification of product structure. Fundamentally, if the material being dried is originally a liquid, then all of those methods that can modify the habit and size of solvent crystals can dramatically alter the rate of sublimation, and hence the time necessary for the primary drying stage. The sublimation rate can be expressed as: 140 5 Intensification of Freeze-Drying for the Pharmaceutical and Food Industries

p p ¼ w;i w;c ; ð : Þ jw 5 10 Rp fl where Rp is an overall resistance to vapor ow, which is the sum of the product resistance, the vial and stopper resistances, and the drying chamber resistance. It has been shown that the product resistance is predominant as it typically contributes more than 90% of the overall resistance (Pikal, 1990). As will be discussed, the above-described resistance is dramatically influenced by freezing conditions and product formulation (see also Chapter 3 in Volume 4, Modern Drying Technology). As shown in Eq. 5.10, in order to increase the sublimation rate, lyophilization professionals can act not only on the product resistance but also on the pressure

difference. The quantity pw,i in Eq. 5.10 is the vapor pressure at the interface of sublimation, which is a function of product temperature. Therefore, in order to maximize the driving force for mass transfer, freeze-drying cycles are designed in

order that product temperature, hence pw,i, is as close as possible to its limit value. A detailed review and comparison between the various methods proposed to date to achieve the above objective is given by Pisano et al. (2013). Rather, in this section attention is focused on the methods proposed to promote the formation of large ice crystals, such as the use of cosolvents or the regulation of nucleation temperature, as these may have a profound impact on product resistance.

5.3.1 Regulation of Nucleation Temperature During Freezing

During freezing, a solvent is separated from the drug and excipients (or nutrients in the case of foodstuffs) by ice formation; the system is then cooled further to allow the transformation of all other components in a frozen state. Various factors affect the freezing characteristics of the solution prior to drying, the most important being the filling volume and the type of container (Hottot et al., 2007), the time–temperature profile (Rey and Bastien, 1961; Searles et al., 2001b), and the supercooling and ice nucleation temperature (Searles et al., 2001a). Control of the above critical variables is essential not only to achieve optimized cycles and to produce uniform batches, but also to simplify cycle scale-up. Lyophilization professionals often focus on primary and secondary drying, but put very little attention to freezing. Yet, freezing is a crucial step for the freeze- drying process as it can impact on both the product characteristics and process performance (Patapoff and Overcashier, 2002; Kasper and Friess, 2011). The product characteristics can be modified only if the nucleation process can be controlled. The control of ice nucleation is further desirable as it forces the nucleation of a batch of vials to occur within narrow temperature and time ranges, and thus enables a more uniform drying behavior to be obtained. In fact, the random nature of the nucleation process causes the nucleation temperature to be distributed around an average value and, as a consequence, vials of the same batch can have different product structures and hence show different behaviors during 5.3 Process Intensification in Vacuum Freeze-Drying of Liquids 141 primary and secondary drying. The stochastic nature of nucleation also makes it difficult to control the temperature within a desired range of supercooling. During the past few decades, however, a variety of methods have been proposed to induce ice nucleation. Inada et al. (2001) suggested the use of ultrasonic vibrations to induce a phase transition in a metastable material. Although the precise mechanism of this phenomenon is unclear (Zhang et al., 2001; Saclier et al., 2010a), various authors have shown that the technique can be used for the effective control of the size of ice crystals and to produce uniform freezing (Saclier et al., 2010b). Unfortunately, however, the ice structure is not uniform within individual vials, and ice crystals tend to be larger at the top surface (Nakagawa et al., 2006). In the “ice fog” technique proposed by Rambhatla et al. (2004), a flow of nitrogen at low temperature and high pressure is introduced inside the drying chamber, when the product has reached the desired value of temperature for nucleation. This operation produces ice crystals that are suspended in the chamber gas: they enter the vials and act as nucleation agents for ice formation. In this way, the nucleation of ice can be induced at a controlled temperature. Unfortunately, the ice structure is usually not uniform in all vials as the time required for nucleation is typically too long. Hence, in order to allow a more rapid and uniform freezing, Patel et al. (2009) proposed that the pressure in the chamber should be reduced when nitrogen is introduced. Currently, the use of this technique is still limited to laboratory-scale freeze-dryers, as large-scale production units may require the use of convection devices to obtain a uniform distribution of the “ice fog” in the drying chamber. Bursac et al. (2009), and more recently Konstantinidis et al. (2011), presented an effective method to control ice nucleation, namely the depressurization method. In this case, vials are loaded onto the freeze-dryer and pressure is increased above atmospheric level by 1.7–2 bar by the bleeding of inert gases at high pressure. The solution is then cooled to a desired temperature and pressure is decreased to induce ice nucleation simultaneously in all vials. Konstantinidis et al. (2011) showed that this method can be applied to both laboratory-scale and production- scale freeze-dryers. If the freeze-dryer can withstand pressures higher than atmospheric value, then the method can easily be implemented; otherwise, it will be necessary to modify the equipment to support overpressure, and this can be expensive. In this situation the application of a different method for ice nucleation control is suggested. “Vacuum-induced surface freezing” is another method of inducing ice nuclea- tion, which can be applied to pre-existing apparatus without any modifications and, thus, may represent a good alternative to the above method. The vials are first loaded onto pre-cooled shelves, after which the pressure is reduced to almost 1 mbar in order to induce water evaporation. As the evaporation is endothermic, this reduces the temperature of the solution at the top surface and promotes its freezing (Kramer et al., 2002). Once nucleation has been induced, the product temperature must be reduced below the melting point and atmospheric pressure restored rapidly in order to prevent the product from puffing up (because of boiling); otherwise, the ice formed at the vial top will melt again (Liu et al., 2005). 142 5 Intensification of Freeze-Drying for the Pharmaceutical and Food Industries

Fig. 5.7 Scanning electron microscopy images of mannitol (metallized samples) that have been freeze-dried from a 5% (w/w) water solution with (a) spontaneous nucleation and (b) forced nucleation. Samples processed in research laboratories of Telstar S.A., Terrassa, Spain.

Figure 5.7 compares the internal structure of freeze-dried products obtained by spontaneous nucleation and by forced nucleation. As expected, the latter sample has larger pores than the sample obtained by spontaneous nucleation. As a result, samples obtained by forced nucleation are characterized by a low product resistance to vapor flow and, therefore, by faster drying rates (see Eq. 5.10). Visual observation confirmed that all vials were nucleated within a short time after pressure decrease and, as a result, they were nucleated at very nearly the same temperature (within 0.1 C) (see Fig. 5.8). The above conditions should guarantee uniform freezing, and this hypothesis was confirmed experimentally (see Fig. 5.9b). In fact, using forced nucleation a

Fig. 5.8 Product temperature profiles during cooling for (a) spontaneous nucleation and (b) forced nucleation of 5% mannitol. Data courtesy of M. Galan and R. Bullich, Telstar S.A., Terrassa, Spain. 5.3 Process Intensification in Vacuum Freeze-Drying of Liquids 143

Fig. 5.9 Comparison of (a) product freeze-drying cycles of 10% mannitol. The temperature as estimated by pressure rise test vertical line indicates the endpoint time for ice technique and (b) chamber pressure profiles sublimation. Data courtesy of M. Galan and R. during primary drying for (, solid line) Bullich, Telstar S.A., Terrassa, Spain. uncontrolled and (, dashed line) controlled sharp change was observed in the Pirani–Baratron pressure ratio curve close to the completion of ice sublimation. In contrast, the cycle with spontaneous sublimation showed a smooth change of the slope for this curve. It should be noted that, although the freezing is uniform, the internal structure of individual vials is heterogeneous, with a compact layer being present at the product top while the underneath part was characterized by much larger pores. A similar structure was observed for samples nucleated by ultrasonic vibrations (Barresi and Fissore, 2011). Figure 5.9 also shows that the drying time in case of forced nucleation was 10 h shorter than the time required when spontaneous nucleation was used, which corresponds to a 40% reduction in drying time. The use of forced nucleation produces large ice crystals (see Fig. 5.7), which leads to a lower product resistance to vapor flowandthusahighrateofsublimation, although drying was carried out at almost the same product temperature (see Fig. 5.9a). As shown in Section 5.5.2, ultrasound technology can also be used to enhance mass transfer during drying. A number of other methods have also been proposed to modify the freezing of liquids in vials. These are aimed at obtaining a uniform freezing, without any concerns regarding the duration of the subsequent drying stages. Although these techniques cannot be used to modify the habit and size of ice crystals, they can induce nucleation in all vials of the batch within a limited temperature range, and this leads to uniform freezing. Searles et al. (2001a) investigated the use of various vial pretreatments (such as roughening, scoring and scratching) as a means of lowering the degree of supercooling, while Petzold and Aguilera (2009) used insoluble impurities as nucleating agents. Unfortunately, none of the above methods allows a precise control of the ice nucleation and, in turn, of the final structure of the freeze-dried product. 144 5 Intensification of Freeze-Drying for the Pharmaceutical and Food Industries

5.3.2 Use of Organic Solvents and Cosolvents

The use of a strictly organic or organic/water system is beneficial to both product quality and process optimization. The potential advantages and disadvantages of using organic solvents and cosolvents (mixed with water, which until now has been the most commonly used solvent) in freeze-drying have been widely discussed by Teagarden and Baker (2002). The main advantage of solvent use, in addition to increasing the solubility of the product, is the increase in the rate of sublimation and hence a decrease in drying time. Before a specific solvent is used in the manufacture of a parenteral product, however, lyophilization professionals must carefully weigh up the associated advantages and disadvantages. The use of organic liquids, either as solvent or cosolvent, produces larger ice crystals, and thus increases the average diameter of the pores created during ice sublimation. As a result, the product resistance to vapor flow is decreased and, if the product temperature does not change, the rate of sublimation will be increased (see Eq. 5.10). Kasraian and DeLuca (1995a) showed, for example, that the product resistance for lactose and sucrose formulations significantly decreased, allowing the sublimation rate to increase, when the product was freeze-dried using a 5% (v/v) aqueous solution of tert-butyl alcohol, instead of pure water. In addition to the reduction in product resistance, most cosolvents used for freeze-drying applications increase the rate of sublimation because they have a higher vapor pressure than water, and hence increase the driving force for mass transfer (Wittaya-Areekul, 1999). A further reduction in energy consumption also results from the fact that the sublimation enthalpy of organic solvents is smaller than that of frozen water. Figure 5.10 compares the product structure obtained by the freeze-drying of 5% (w/w) sucrose solution with pure water and with a 5% (w/w) aqueous solution of tert-butanol. These images show clearly that a different structure (with needle-like crystals) is obtained when 5% tert-butanol is used compared to water. Figure 5.11 compares two freeze-drying cycles for a 5% (w/w) aqueous solution of sucrose; the former was carried out using pure water as solvent,

Fig. 5.10 Scanning electron microscopy images of sucrose samples (metallized) that have been freeze-driedfroma5%(w/w)watersolutionwith(a)purewaterand(b)5%(w/w)aqueoussolutionof tert-butyl alcohol. 5.3 Process Intensification in Vacuum Freeze-Drying of Liquids 145

Fig. 5.11 Example of freeze-drying cycles for a Evolution of (a) product temperature at vial 5% (w/w) sucrose solution with (solid line) pure bottom as detected by thermocouples and water and (dashed line) 5% (w/w) tert-butyl (b) water concentration in chamber gas as alcohol aqueous solution. Primary drying was measured by a laser detector. The vertical line carried out at Tshelf ¼ –20 C and Pc ¼ 10 Pa. indicates the endpoint time for ice sublimation.

while in the latter case a 5% (w/w) aqueous solution of tert-butyl alcohol was used. The same processing conditions (i.e., temperature of the heat transfer fluid and chamber pressure) were used for both cycles. The drying rate was significantly increased when tert-butyl alcohol was used as cosolvent, and consequently the drying time was also decreased by approximately 50%. In fact, as shown in Fig. 5.11b, the water concentration in the drying chamber began to decrease when ice sublimation was complete – that is, after 12.5 h for the cycle with pure water and after 23.5 h for 5% (w/w) tert-butyl alcohol solution. The increase in drying rate was most likely due to the formation of needle-shaped ice crystals (see Fig. 5.10b), which dramatically lowered the product resistance to vapor flow. It should also be noted that product temperature was slightly lower than the value observed when sucrose was freeze-dried in the presence of tert-butyl alcohol (see Fig. 5.11a). Similar results were shown by Kasraian and DeLuca (1995b). It follows that the product could be successfully freeze-dried in thepresenceoftert-butyl alcohol by using a more aggressive heating policy than that used for aqueous solutions. As a result, a further reduction could be obtained for the drying time when a cosolvent system is used. Of course, the above considerations are valid only if the presence of tert-butyl alcohol does not lower the collapse temperature of the formulation. For sucrose- based formulations, Kasraian and DeLuca (1995a) showed that the addition of tert- butyl alcohol does not alter the collapse temperature; however, there are cases where the presence of an organic solvent together with water can reduce the collapse temperature of the formulation. As an example, Liu et al. (2005) showed that the collapse temperature of sulfobutylether-7-b-cyclodextrin formulation was reduced by 2 C when a 5% (w/w) tert-butyl alcohol aqueous solution was used. 146 5 Intensification of Freeze-Drying for the Pharmaceutical and Food Industries

Tesconi et al. (1999) proposed the use of organic solvents that are solid at room temperature as a means of freeze-drying pharmaceutical products, without using conventional freeze-drying equipment. These authors investigated a variety of compounds in terms of their ability to solubilize hydrophobic drugs, to increase the chemical stability of both the pre-dried bulk solution and the dried product, to be readily removed by vacuum drying, and to produce elegant cakes which can easily be reconstituted. Ultimately, an eutectic mixture of chlorobutanol hemihydrate and dimethyl sulfone was found to be an optimum candidate, and the use of this solvent made it possible to carry out the freeze-drying process for a drug model without freezing. Furthermore, the process required only modest heating under vacuum. Unfortunately, this method did not find real application in industry due to various limitations; for example, solutions must be kept in the liquid phase during filtration and filling operations, while a high degree of flammability can cause problems of operator safety. Nonetheless, the application of this method to production units would represent a further step towards process intensification for freeze-drying, though a more accurate study to determine its technical feasibility may be necessary.

5.4 Atmospheric Freeze-Drying

Vacuum freeze-drying produces dried products that retain almost all of their original characteristics, such as color, flavor, and taste. Moreover, the high specific surface area generally allows an easy and rapid rehydration. The main drawback of the process relates to the cost of the operation, as fixed costs can be high due to vacuum requirements while energy costs can be significantly higher compared to other drying processes (the specific moisture extraction rate in a vacuum freeze- drying process is in the range of 0.4 kg of water per kWh, as reported by Claussen et al., 2007c). In order to reduce the energy consumption of the process, and thus to improve its sustainability, atmospheric freeze-drying (i.e., drying with cold air or nitrogen at normal pressure) has been proposed, in which case it would be possible to achieve a specific moisture extraction rate ranging from 1.5 to 4.6 kg of water per kWh (Claussen et al., 2007c). In fact, Wolff and Gibert (1990a) reported that an up to 35% saving of energy costs was achievable by using atmospheric freeze-drying instead of vacuum freeze-drying for potato slices. Meryman (1959) first showed that the pressure gradient in the dried cake, rather than the absolute pressure in the system, determined the vapor flow from the interface of sublimation to the drying chamber. This permits a product to be freeze- dried at atmospheric pressure, provided that a vapor pressure gradient is maintained in the dried product. The atmospheric freeze-drying of various products has been investigated previously, examples being fish products and starter feeds for the fish industry, strawberries and potatoes and lactic acid bacteria, among many others (Alves-Filho, 2002; Alves-Filho et al., 2004; Alves-Filho and Roos, 2006; Alves-Filho and Eikevik, 2008). Attention must be paid, however, to 5.4 Atmospheric Freeze-Drying 147 avoid any possible confusion between atmospheric freeze-drying and the drying of a frozen product at atmospheric pressure. In fact, for some reported applications of atmospheric freeze-drying the processing conditions used are so aggressive that they instill a doubt that the temperature of the initially frozen product will increase up to the melting temperature of ice. But, if this is the case, ice will melt during the process, the product to be dried will no longer be frozen, and the term “freeze- drying” would not be appropriate to describe such a drying process. In fact, an atmospheric freeze-drying process consists of the convective drying of a product maintained at a temperature below its freezing point: typically, the temperature ranges from 10 to 3 C, as this appears to be a good compromise between final product quality and costs, as lower temperatures reduce the possibility of removing moisture with air and thus increase the duration of the process. The possibility of reducing drying time by using a variable process temperature has also been investigated. Indeed, it has been noted that an increasing temperature can inhibit solute transport within the product and enhance the heat transport in the dried product, although in some cases a lower temperature and a longer residence time can improve the product quality (Strømmen et al., 2005). The majority of published reports have dealt with atmospheric freeze-drying in fluidized-bed dryers (Wolff and Gibert, 1990a; Lombrana~ and Villaran, 1996) and also in spray freeze-dryers (Mumenthaler and Leuenberger, 1991; Leuenberger, 2002; Rogers et al., 2003; Leuenberger et al., 2006). When atmospheric freeze- drying is carried out in a fluidized bed, we can take advantage of the high values of the heat and mass transfer coefficients; notably, the product must be frozen and granulated before drying. One drawback of this process, however, is represented by the size reduction caused by mechanical cracking, and as an alternative it is possible to carry out the process in a tunnel dryer, even if the heat and mass transfer is worse (Claussen et al., 2007b). Spray freeze-drying appears to be a valuable alternative to produce a free-flowing powder, with a high surface area, a porous end product, and good instant characteristics, with enhanced solubility and a uniform and ultrafine particle size. Spray freeze-drying into liquids, gases (e.g., a refrigerated air stream), and into gases over a fluidized bed have been reported. Recent investigations have been focused on the inclusion of a heat pump in the drying system (Indergard et al., 2001; Alves-Filho, 2002; Chua et al., 2002), as this not only allows significant energy savings but also permits a careful control of the drying temperature and air humidity. A comprehensive treatment of heat pump drying was provided in Chapter 4 of Volume 4, Modern Drying Technology. The kinetics of atmospheric freeze-drying can be investigated by means of mathematical modeling, and the receding ice front model is perhaps the most common type of model used for this purpose. In this case, the product is considered to be divided into two layers, namely the frozen inner core and the dried outer layer. A receding interface is assumed to exist between the frozen and the dried layer. Subsequently, heat transferred from the surface of the product to the interface is used for sublimation of the ice, which diffuses from the interface to the surface and to the drying chamber at atmospheric pressure. In the case of a 148 5 Intensification of Freeze-Drying for the Pharmaceutical and Food Industries

T p air w,air drying chamber

  dried product T Q M p i w,i interface of wet product sublimation

Fig. 5.12 Sketch of an atmospheric freeze-drying process.

layered material with a planar sublimation front (Fig. 5.12), the transport of heat and mass through the dried layer can be modeled by means of the following equations (Wolff and Gibert, 1990b; Claussen et al., 2007a): _ A Q ¼ ðÞTair Ti ; ð5:11Þ 1 þ Ld aext ld 1 _ ¼ A ; ð5:12Þ M ~ 1 pw;i pw;air RT i þ Ld bext D a b fi where ext and ext are, respectively, the heat and the mass transfer coef cient from the external surface of the product to the gaseous stream, and D is the diffusion coefficient. As the heat flux arriving at the interface between the frozen and the dried product is used for ice sublimation, then: Q_ M_ ¼ : ð5:13Þ Dhs Equations 5.11–5.13 can be used to calculate the interface temperature, once the operating conditions (temperature and humidity of the air stream) have been selected, and the values of the heat and mass transfer coefficients have been determined (see, among others, Boeh-Ocansej, 1988; Krokida et al., 2002). It is then possible to calculate the sublimation flux and, finally, the variation of the dried layer thickness with time: dL Ar ðÞX X d ¼ M_ : ð5:14Þ d 0 fin dt Figure 5.13 shows the effect of air velocity, temperature, and relative humidity on drying time for the atmospheric freeze-drying of apple samples (Malus domestica) in a tunnel dryer. It appears that, while the drying time is only slightly reduced when the air velocity is strongly increased, the effect of the air temperature and of the relative humidity is much more relevant, and the drying time can be greatly reduced by increasing air temperature and decreasing the relative humidity. 5.4 Atmospheric Freeze-Drying 149

Fig. 5.13 Effect of (a) air velocity (air relative thickness ¼ 5 mm, D ¼ 5 10 5 m2 s 1)ina humidity ¼ 30%) and (b) air relative humidity tunnel dryer for different values of air (air velocity ¼ 3ms 1) on the drying time of temperature (& ¼ 0 C, ¼5 C, apple (Malus domestica) samples (product ~ ¼10 C, ^ ¼15 C).

In case the interface of sublimation is not planar, as for particles in fluidized-bed drying, the heat and mass balance equations must be modified to account for this (Wolff and Gibert, 1990b; Di Matteo et al., 2003). When comparing the calculated drying time for a tunnel freeze-dryer and a fluidized bed, as shown in Fig. 5.14 for the atmospheric freeze-drying of apple samples (M. domestica), it appears that the drying time is highly reduced in the second type of equipment (due mainly to the higher heat and mass transfer coefficients). Lower values of relative humidity and higher values of temperature result in this case in a decreasing drying time.

(a) (b) 60 60

50 50 , h , h 40 40

30 30

20 20 drying time drying time

10 10

0 0 %1 10% 30% 1% 10% %03 air relative humidity air relative humidity

Fig. 5.14 Comparison between the drying (Malus domestica) samples as a function of air time calculated for a tunnel dryer (&,air relative humidity and temperature (a) ¼ 1 ¼ ¼ ¼ velocity 3ms , product thickness 5 mm) Tair 5 C; (b) Tair 10 C. and in a fluidized-bed dryer ( ) for apple 150 5 Intensification of Freeze-Drying for the Pharmaceutical and Food Industries

5.5 Use of Combined Technologies for Drying Heat-Sensitive Products

As discussed above, atmospheric freeze-drying produces high-quality products but still entails long drying times. In the past, various research groups have striven to decrease energy consumption in atmospheric freeze-drying, for example, by new technologies that use alternative forms and sources of energy for processing. These technologies can be applied to either enhance heat transfer between product and heat source, such as microwave, radiofrequency and infrared radiation, or simply to intensify the rate of dehydration without increasing the amount of heat supplied to the product, for example, by using high-intensity sonic and ultrasonic waves. The advantages and disadvantages of the above techniques are discussed in the following subsections, with emphasis placed on their application to production equipment.

5.5.1 Microwave-Assisted Drying

When microwaves are used as an energy source, heat is generated within the product being dried. This energy is transferred by molecular excitation (i.e., rotation of molecules with permanent dipole moment and ionic motion) caused by the alternating electromagnetic field and, as a result, heat is supplied to the entire mass of the wet product in a rapid and uniform fashion. The magnitude of the above effect is product-specific as it depends on the dielectric properties of materials being heated, and therefore its application in drying must be evaluated for each individual product. At this point, the use of microwaves is first presented as an alternative energy source to be combined with various drying methods, such as hot-air drying, vacuum-drying, and heat-pump drying. Attention is then focused on the applica- tion of microwave energy to the freeze-drying process, evidencing its advantages and current limitations. Tulasidas et al. (1995) showed that the application of microwave energy to drying processes constitutes an efficient method of transferring heat to products for water removal. As the use of microwaves alone cannot complete a drying process, this technique commonly is combined with forced air streams and vacuum conditions (Chou and Chua, 2001). A number of studies have shown the benefits of using microwave-assisted drying for vegetables; a list of recent applications is given in Tab. 5.1, where microwave technology is used to support both hot-air drying and vacuum-drying. In particular, various groups have shown that microwave-assisted vacuum drying can produce high-quality dried products, which often can be rated as equal to freeze-dried products. Kwok et al. (2004) compared various drying techniques (including freeze-drying) for preserving the antioxidant properties of Saskatoon berries. The relatively low consumption of energy, together with a significant reduction of drying time, observed for microwave-assisted technology has led to it being competitive with 5.5 Use of Combined Technologies for Drying Heat-Sensitive Products 151

Tab. 5.1 Examples of microwave-assisted drying for foodstuffs.

Material Microwave-assisted Reference technology

Banana Vacuum drying Drouzas and Schubert, 1996 Carrot Convective drying Wang and Xi, 2005 Vacuum drying Lin et al., 1998 Regier et al., 2005 Convective/vacuum Orsat et al., 2007 drying/ Cabbage Vacuum drying Xu et al., 2004 Cranberry Vacuum drying Yongsawatdigul and Gunasekaran, 1996a, 1996b Garlic Convective drying Sharma and Prasad, 2006 Convective/vacuum drying Sunjka et al., 2004 Vacuum drying Cui et al., 2003 Ginseng Vacuum drying Popovich et al., 2005 Mushroom Convective drying Torringa et al., 2001 Pea Atmospheric freeze-drying Eikevik et al., 2012 Raisin Convective drying Tulasidas et al., 1995 Saskatoon Vacuum drying Kwok et al., 2004 berry Strawberry Vacuum drying Venkatachalapathy and Raghavan, 1999 Bohm€ et al., 2006

freeze-drying, even if the former can entail a potential loss of activity. The best retention was obtained by freeze-drying followed by microwave-assisted vacuum drying. Similar results were obtained for the drying of garlic cloves by Cui et al. (2003). Sharma and Prasad (2006) also observed that microwave-assisted drying allowed an 80% reduction in drying time when compared to hot-air drying, while an even larger reduction can be obtained if compared to freeze-drying process. When Lin et al. (1998) compared freeze-drying, hot-air drying, and microwave- assisted vacuum drying for carrot slices, they observed that both freeze-drying and microwave-assisted vacuum drying can produce a high-quality dried product. Later, Regier et al. (2005) showed that microwave-assisted vacuum drying could produce dried carrots with the highest carotenoid retention when compared to both freeze- drying and hot-air drying. Popovich et al. (2005) compared the type and content of ginsenosides contained in ginseng that had been dried using various methods, and found freeze-dried ginseng to have the highest contents of Rg1 and Re ginsenosides. In contrast, microwave-assisted vacuum drying produced a dried product with the highest contents of Rb1 and Rd ginsenosides. Finally, Venkatachalapathy and Raghavan (1999) observed that a combination of microwave-assisted vacuum drying and osmotic dehydration (as pretreat- ment), when applied to the drying of strawberries, created a high-quality dried 152 5 Intensification of Freeze-Drying for the Pharmaceutical and Food Industries

productwhichwascomparablewiththatobtainedbyfreeze-drying,butina much shorter time. Various groups (e.g., Ma and Peltre, 1975; Ang et al., 1977) have shown that microwave technology can also be applied to the freeze-drying of foodstuffs, with a dramatic increase in sublimation rate. Eikevik et al. (2012) showed that the drying time for green peas could be halved when atmospheric freeze-drying was combined with microwave radiation. However, the rate of drying and product quality were dramatically affected by the process conditions, for example, fixed versus fluid-bed conditions. Arsem and Ma (1990) also investigated the possibility of using microwave combined with radiative heat transfer as a means of further increasing the drying rate. The above studies demonstrate that the microwave-assisted vacuum drying of foodstuffs is an effective alternative to freeze-drying, as it produces high-quality dried products in a shorter time and entails a low energy consumption. By contrast, the application of microwave technology to freeze-drying can be problematic, as corona discharge and nonuniform heating can occur and produce high-energy zones (Pester et al., 1977). This phenomenon represents a strong limitation for the drying of pharmaceutical products, which must be dried at very low temperatures; moreover, even small and brief variations in drying temperature can cause damage to the active pharmaceutical ingredients. However, it has been shown that microwaves can be applied successfully to the freeze-drying of food liquids, such as milk foam (Sochanski et al., 1990), as their critical quality attributes are less stringent than those required by the pharmaceutical industry. In conclusion, the use of microwaves as an additional heat source during freeze-drying may be beneficial to process performance, as it can enhance and make more uniform the heat transfer to the product. However, in order to promote microwave application as a potential and effective technology for intensifying the freeze-drying process, further studies are required to improve the control of the energy supplied by microwaves.

5.5.2 Ultrasound-Assisted Drying

The use of ultrasound to intensify drying processes was first investigated by Fairbank (1975), although others (Gallego-Juarez et al., 1999; Mulet et al., 2003) have more recently applied direct contact ultrasound to the drying of vegetables and shown the technique to allow dramatic reductions in drying times (cf. Chapter 8 of this Volume). Unlike previous technologies, the use of ultrasound does not constitute an alternative source of heat for processing (increase in product temperature is significant only if the material is exposed for a long time to ultrasound) and, therefore, does not introduce concerns about product overheating or thawing. Various groups have shown that power ultrasound increases the rate of water removal based on mechanical actions on both gas–solid interfaces (Bantle and Eikevik, 2011) and product structure (Garcia-Perez et al., 2012b). The application of high-intensity ultrasonic energy produces rapid contractions and 5.5 Use of Combined Technologies for Drying Heat-Sensitive Products 153 expansions through the product structure, which in turn promotes the formation of microscopic channels for water motion. High-intensity ultrasonic waves also induce cavitation, which enhances the removal of water that is strongly linked to the solid product (Muralidhara et al., 1985). The use of high-intensity ultrasound technology, combined with atmospheric freeze-drying, has recently been investigated by Bantle and Eikevik (2011) for the drying of peas in fluidized beds. In particular, the use of ultrasound was shown to increase the effective diffusivity by up to 15%, and that such an increase was higher at low temperatures (between 6 Cand0C) but was marginal at room temperature. A similar result was obtained by Garcia-Perez et al. (2012a), who showed that the use of ultrasound in the convective drying (at atmospheric pressure and low temperatures) of vegetables could reduce the drying time by up to 70%. Such a reduction in drying time was associated with a dramatic increase in vapor diffusivity of up to 400%. Ultrasound technology was also shown to speed up the drying process when an organic solvent (i.e., ethanol) was used in place of water. Figure 5.15 compares the structure of albedo cells as obtained from orange peel in the case of a product dried with or without ultrasound technology. It is well known that drying produces high stress on vegetable cells and can lead to a partial collapse of the material structure (Garcia-Perez et al., 2012b). This phenomenon allows a reduction in the product resistance to vapor flow and, hence, speeds up the dehydration process. A similar behavior was also observed when drying was combined with ultrasound technology although, as shown in Fig. 5.15, in this case the product damage that occurred during drying was more marked. Because of this phenomenon, the product resistance to vapor flow can be further reduced and the rate of water removal increased. The results of the above-described studies have demonstrated that ultrasound technology can be effectively coupled with low-temperature drying to speed up water removal and, hence, lead to cost-effective processes. Nevertheless, the use of this technique in production units remains limited, most likely because of practical difficulties in satisfying industrial-scale requirements.

Fig. 5.15 Scanningelectronmicroscopy images of albedocells obtained from orangepeel in case of (a) hot-air drying and (b) ultrasound-assisted hot-air drying. Illustration courtesy of Prof. Antonio Mulet and Juan A. Carcel, Universidad Politecnica de Valencia, Spain. 154 5 Intensification of Freeze-Drying for the Pharmaceutical and Food Industries

5.6 Continuous Freeze-Drying

During the 1960s, the food industry was forced to increase its production to cope with increased requests from the market for freeze-dried foodstuffs, such as milk and coffee. Surprisingly, the above-described objective was achieved not by improving preexistent equipment but, quite oppositely, by abandoning it in favor of new technologies. The food industry, in fact, revamped its equipment and moved away from batch to semi-continuous and, as a last resort, to continuous operations. A first attempt at continuous operation was developed during the 1960s by Leybold and Oetjen, who created the first prototype continuous freeze-dryer for milk products (Oetjen and Haseley, 2004; Rey, 2010). The operation encompassed several steps. The product was first frozen to produce granules that were then loaded onto trays which hang on a monorail. In this way, the product could be moved through the freeze-drying tunnel between temperature-controlled shelves, while the tunnel was divided into several sections, separated by sliding gates. This approach allowed a dramatic reduction in the total cycle time, namely hours versus days. In an alternative configuration, the frozen products were loaded tray-by-tray through a small side lock. The above technology was used to freeze-dry not only coffee and milk but also fish fillets, meat, and vegetables. Although the above-described technology allowed a dramatic reduction in energy consumption and processing costs, as well as a desired increase in the rate of production, the process was semi-continuous and could not satisfy the increased requests of the international market for instant coffee. In fact, this request motivated the food industry to develop a fully continuous line, which was realized in the mid-1970s. Continuous plants for freeze-drying are highly sophisticated and were designed by the food industry itself, which kept the development confidential. In one of the most advanced configurations, granules of frozen product were continuously deposited onto long vibrating trays through a rotating lock. The product being dried moved on trays as a fluidized bed, such that the granules were forced to move into parallel channels through vertical ribs. During the plant design, great attention was paid to the problem of dust formation; for this purpose, various solutions were investigated to prevent granule-to-granule attrition, as well as attrition between the granules and tray walls. Although such attrition was limited, dust formation proved to be unavoidable, and semi-cylindrical screens were installed over the vibrating trays in order to trap any dust produced by mechanical shocks; this prevented the dust from being transported to the vacuum unit (i.e., condenser and vacuum pumps) via the vapor stream. In fact, the use of continuous freeze-drying led to an enormous increase in process efficiency to a point where the total cycle time was of the order of minutes (compared to a few hours for a semi- continuous plant, and days for batch units), throughput could reach tens of tons per day, and the dried products were of a constant and uniform quality. Although the freeze-drying process has been greatly intensified in the food industry by switching from batch to semi-continuous and finally to continuous plants, biological and pharmaceutical products are still largely produced in batch 5.7 Conclusions 155 units. Typically, the low efficiency of these processes was justified by the high value of the end products, but today this situation has changed as market competition continues to grow, mainly for those products which are no longer intellectually protected. In addition, manufacturers still face recurrent problem of heterogeneity in end product quality within a batch. Although various solutions were investigated and successfully applied in order to contain this latter problem, waste due to product heterogeneity remains high. Furthermore, validation tests proposed by regulatory agencies are increasingly stringent. Today, with multi-shelf freeze-dying cabinets capable of being loaded with up to 100 000 vials, it is unthinkable that vials in different positions can be dried in the same way, even if elaborate designs were proposed by the freeze-dryer manufacturers. One possible solution to this problem would be to switch to a semi-continuous process, as has the food industry during the past 50 years (Rey, 2010), whereby the freeze-drying process can be greatly intensified in terms of throughput, energy consumption, and product quality.

5.7 Conclusions

In freeze-drying, process intensification involves the development of new technologies that can reduce consumption costs and lead the way towards uniformity in end-product quality. The former objective can be achieved by reducing the drying time. As primary drying is the longest and the most energy-consuming stage, its optimization is typically the primary objective, while minor attention is dedicated to the optimization of secondary drying. Various solutions have been discussed in this chapter to achieve the above-described objectives. An example is the use of automatic control to optimize processing conditions; that is, to find in-line a combination of pressure and temperature that maximizes the rate of sublimation, while satisfying constraints posed by the equipment, as well as by product quality attributes (e.g., on the maximum value for product temperature). The control of freezing can also benefit process intensification, as this stage can impact the product structure and hence modify the product resistance to mass transfer. Various groups have also claimed that it can improve batch uniformity. In this chapter, the various methods proposed to date to regulate ice nucleation have been briefly described, and an example of application to a real case provided. An alternative approach to enhancing the drying rate is to use organic solvents, the presence of which leads to larger solvent crystals and hence a lower product resistance to vapor flow. In addition, organic solvents are typically characterized by high values of vapor pressure and so can increase the driving force for mass transfer. Both of these phenomena produce high rates of sublimation. Atmospheric freeze-drying represents a good alternative to vacuum freeze- drying, as a means of increasing process efficiency. Further improvements can be obtained by combining atmospheric or vacuum freeze-drying with new technolo- gies that can then be used to enhance heat transfer between the heat source and the 156 5 Intensification of Freeze-Drying for the Pharmaceutical and Food Industries

product (microwave-assisted drying), or simply to intensify the rate of dehydration by altering the product structure (e.g., high intensity sonic and ultrasonic waves). The application of microwave technology to pharmaceutical products can be problematic, however, as the latter products must be dried at very low temperatures and the risk of product overheating is high. In contrast, practical difficulties to satisfy industrial-scale requirements limit the application of high-ultrasound techniques. A further step towards process intensification is provided by continuous plants, and such technologies have been used for food manufacture since the 1960s, allowing dramatic increases in throughput and product quality uniformity. In theory, continuous plants may also be developed for the manufacture of pharmaceuticals, although regulatory concerns may still impede this pathway.

Acknowledgments

The authors acknowledge Irene Oddone, Susanna Sanapo, Salvatore Genco and Alessandro Ando for their contributions to this chapter. The contributions of Miquel Galan and Robert Bullich (Telstar s.a., Terrassa, Spain), and of Prof. Antonio Mulet and Juan A. Carcel, (Universidad Politecnica de Valencia, Valencia, Spain) are also gratefully acknowledged.

Additional Notation Used in Chapter 5

fi 1 1 cp speci c heat capacity at constant pressure J kg K Ex exergy J L layer thickness m Q energy J fl 1 Rp overall resistance to vapor ow m s

Subscripts and Superscripts

c drying chamber cond condenser cooling cooling d drying, dry des desublimation ext external fin final freeze, f product after freezing, frozen References 157 h heat transfer i interface of sublimation ice ice in initial, input, inlet value loss loss m mass transfer out outlet pd primary drying pump pump s sublimation shelf heating shelf v vapor 0 initial, reference

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6 Drying of Foamed Materials Ireneusz Zbicinski, Julia Rabaeva, and Artur Lewandowski

6.1 Introduction

A large group of thermally sensitive materials exists, such as mixtures of biological substances, proteins, fats, vitamins and mineral salts, which are difficult to dry by using classical techniques. One method by which to overcome this problem is to foam the material before drying, and two basic techniques of foamed material drying have been developed: (i) the drying of dispersed foams (spray drying); and (ii) the drying of bulk foams (foam-mat drying). A conventional spray drying process is modified to foam spray drying by the injection of an inert gas (e.g., N2,NO2,CO2)intothefeedstockwiththe following formation of foamed droplets during atomization. Compared to the standard spray drying process, the addition of a foaming gas increases the dryer throughput, intensifies the thermal process efficiency, reduces the thermal degradation of the product and increases the retention of highly volatile sub- stances (Hanrahan and Webb, 1961; Bell et al., 1963; Abdul-Rahman et al., 1971; Crosby and Weyl, 1977). Investigations carried out over the past few decades have revealed that the foam spray drying process also affects the final product properties, such as particle diameter, porosity, and bulk density (Hanrahan and Webb, 1961; Hanrahan et al., 1962). In the foam-mat drying method, a liquid raw material is first whipped to produce a bulk volume of foam which is then air-, freeze-, vacuum-, or microwave-dried. Selected surfactants relevant to the properties of the product must be added in order to obtain a stable foam. Foam-mat drying, which is mainly applied to dry liquid or semi-liquid foods, offers a reduction in drying time at low drying temperatures due to the increased interfacial area of foamed materials, a high drying efficiency, and a quick rehydration of the product. The main drawbacks of the foam-mat drying process include foam stability problems, low dryer through- put, and a loss of aromatic components (Marques et al., 2006). Numerous reports have also been made describing a process similar to foaming denoted by aerification or aeration, in which air is introduced to the feed in order to increase the oxygen concentration or to change the rheological properties (Zuniga

Modern Drying Technology Volume 5: Process Intensification, First Edition. Edited by Evangelos Tsotsas and Arun S. Mujumdar. Ó 2014 Wiley-VCH Verlag GmbH & Co. KGaA. Published 2014 by Wiley-VCH Verlag GmbH & Co. KGaA. 164 6 Drying of Foamed Materials

and Aguilera, 2008; Campbell and Mougeot, 1999). Aeration, achieved by bubbling air through the liquid, by spraying the liquid into the air, or by agitation of the liquid is widely used in food manufacturing processes (e.g., ice-cream, bread, cakes, aerated water or cola) and in the chemical industry. An overview of the principles and applications of foam-mat and foam spray drying processes is provided in the following subsections.

6.2 Foam Properties

Foams are disperse systems in which the liquid is a dispersing phase and the gas is a dispersed component. In order to obtain a solution of the consistency of foam, a two-phase liquid–gas system with the application of an appropriate surfactant must be formed (Malysa and Lunkenheimer, 2008). Due to the energy supplied to disperse the gas phase, gas bubbles surrounded by a thin liquid film containing surfactant molecules are released. In the case of pure solutions, the process of foam formation is very difficult because gas bubbles contained in the liquid or solution cannot be stabilized without a surfactant (Prud’Homme and Khan, 1996). A characteristic feature of all foams is low density. Energy must be supplied to the system to obtain a material with the consistency of a foam, both in liquid and solid form, irrespective of its composition and use. Foams can be obtained by both condensation and dispersion methods (Ekserova and Kruglyakov, 1998):

 Condensation methods include all processes in which gas bubbles are formed due to a decrease in the external pressure (boiling, cavitation, desorption, chemical reaction).  In dispersion methods (sieving, gas admixing, mixing), the process of foam formation is connected with gas dispersion which involves the introduction of gas bubbles into the solution.

Bubbling, pneumatic processing or mechanical action (e.g., stirring) are used to

produce foam. Usually, N2,CO2 or atmospheric air, all of which are relatively cheap foam-forming gases, are used in the process (Farajzadeh et al., 2011). Foam is modified during foaming, and the geometric shape of the foam cells, in relation to the foam microstructure, depends on the gas-to-liquid volume ratio. The gas volume in single foam cells can reach even 75% of foam volume. At the first stage of foam formation, a single layer of bubbles forms a thin foam film on the liquid surface. An increase of the gas volume in the liquid then causes foam to further build-up until there is a loss of stability due to the excessively increased volume of gas bubbles (Ekserova and Kruglyakov, 1998). One method of improving foam stability is to increase the concentration of surfactants (e.g., soy, whey, milk protein) and the thickness of the films (Philips et al., 1987). The stability of a foam structure depends on the location of cells and their mutual position in space. 6.2 Foam Properties 165

Fig. 6.1 Differentstructuresofdropletsobtained byatomization.(a)Non-foamed;(b)Foamedwith one bubble; (c) Foamed with many bubbles.

In spray drying, a gas-admixing technique is used to produce a dispersed foam from the feed, whereas in foam-mat drying mechanical methods (e.g., mixing, whipping) are applied to stabilize a bulk volume of the foamed raw material. In foam spray drying, depending on the drying and spraying process parameters, irregular and complex structures of foamed particles can be obtained. Figure 6.1 shows an example of foamed (Fig. 6.1b, c) and non-foamed (Fig. 6.1a) particles obtained as a result of atomization (Rabaeva, 2012). Spray foaming develops the following forms of droplet/particle structure: droplets with no gas inside (Fig. 6.1a); droplets with gas inside (Fig. 6.1b); and droplets with entrapped gas bubbles with inner bridges of liquid or dried solution (Fig. 6.1c). The drying mechanism of a droplet depends on the droplet structure. Abdul- Rahman et al. (1971) and Crosby and Weyl (1977) estimated that the heat and mass transfer rates for foamed droplets (Fig. 6.1b and c) were faster than for non-foamed droplets (Fig. 6.1a), due to an accelerated transport of liquid water to the evaporation front (Ratti and Kudra, 2006) and an increase in the droplet surface area caused by gas expansion during drying. During the foam spray drying process, particles with all of the above structures are formed simulta- neously (Fig. 6.1a–c). Figure 6.2a and b show the structure of foamed and non-foamed spray-dried maltodextrin powder at drying temperature 200 C, respectively (Rabaeva, 2012). On analyzing Fig. 6.2b, particle structures can be observed with voids and inner bridges, which are similar to the foamed droplet structure shown in Fig. 6.1c. The effect of foaming on dried particle morphology was also analyzed by Frey and King (1986), who found that gas-admixing foam spray drying produced particles with many small internal voids that may result from the entrainment of air during the atomization or desorption of dissolved air (Greenwald and King, 1981). Large single internal voids were observed for particles produced by the gas- desorption method. These observations were in agreement with results produced by Hanrahan et al. (1962) and Abdul-Rahman et al. (1971), who investigated particle morphology for whole-milk and sodium caseinate solutions. In the foam-mat drying method liquid foods are whipped into stable foams by using mainly mechanical methods, such as mixing to produce a bulk volume of 166 6 Drying of Foamed Materials

Fig. 6.2 Microscope images of spray-dried powder (maltodextrin, GMS, drying temperature 200 C). (a) Non-foamed; (b) Foamed.

foam. In order to obtain a stable foam it is necessary to add carefully selected surfactants that are relevant to the properties of the product. An example of the bulk volume structure of foam-dried particles (e.g., maltodextrin/sodium caseinate powder) is shown in Fig. 6.3 (Schoonman et al., 2001). Here, the solid matrix, voids, open and closed pores and bubbles, micropores and cracks create a complex structure that affects both heat and mass transfer during drying. Bulk foams processed by the direct foaming of slurries have a tendency to develop defects such as cracks or cell size gradients as a result of a rise in capillary stress due to water evaporation during drying. Today, a vast literature is available on the reduction of microstructural inhomogeneities of foams by means of, mainly, the selection of suitable surfactants, the concentration of surfactants, and the drying method or adjustment of drying process parameters (Whang et al., 1995; Schoonman et al., 2001; Kim et al., 2004; Pisal et al., 2006; Chino and Dunand, 2008; Sundaram and Durance, 2008; Pradhan and Bhargava, 2008; Loa et al., 2011).

Fig. 6.3 Bulk volume of foam-dried particles (maltodextrin/sodium caseinate powder), according to Schoonman et al. (2001); 1: Solid matrix; 2: Voids; 3: Open pores; 4: Closed pores; 5: Cracks; 6: Connected pores. 6.3 Foam Spray Drying 167

6.3 Foam Spray Drying

6.3.1 Processing Principles

The foam spray drying process was developed for dairy industry at the start of the twentieth century. The spray drying of foamed materials was patented in 1917 (Campbell, 1917) to dry foamed milk and egg albumin (Ratti and Kudra, 2006), but some 40 years later Sinnamon et al. (1957) found that the vacuum drying of concentrated milk foam would produce a whole-milk powder of good flavor and dispersibility. Morgan et al. (1959) showed that the dispersibility of whole-milk powder could also be maintained in foamed products during drying at atmospheric pressure. Although, since that time, very few studies on foam spray drying have been reported, this technology has more recently attracted attention because of the possibilities of processing hard-to-dry food materials and to control the final product properties such as bulk density, porosity, solubility, and wettability. Two main methods have been developed for feed foaming applied to spray drying, namely gas-desorption and gas-admixing. In the gas-desorption method, the gas (CO2) is dissolved in the liquid feed under moderate pressure; subsequently, after atomization, droplets supersaturated with the gas are formed, such that the bubbles become nucleated and grow within the droplets. In the gas-admixing method, foaming gas is directly added to the feed. However, in order to obtain a foamed feed, a precisely determined amount of inert gas (e.g., air, N2,NO2,CO2) must be injected into the feed at a high pressure, between the pump and the atomizer. The most important part of the foaming system is a mixing device; typical gas- mixing devices, as applied by Hanrahan and Webb (1961) and Zbicinski and Rabaeva (2010), are shown in Fig. 6.4. A gas-mixing chamber consists of a main pipe through which the feed flows, and a separate channel for the foaming gas injection. After flowing through the gas- mixing chamber, a two-phase feed (solution and gas) is delivered to a mechanical homogenizer to produce fine foaming gas bubbles in the liquid, which is then transferred to the nozzle. An example of a typical set-up of the gas-admixing foaming system for spray drying is shown in Fig. 6.5 (Rabaeva, 2012). The gas-admixing foaming system consists of the following elements:

 tank/mixer of the solution (1)  mono pump (2)  tank with foaming gas, here N2 (3)  gas pressure valve (4)  gas heater (5)  foaming gas flow meter/control (6) 168 6 Drying of Foamed Materials

Fig. 6.4 Gas mixing devices. (a) Adapted from Hanrahan and Webb (1961); (b) Adapted from Zbicinski and Rabaeva (2010).

 foaming (mixing) device (a system for gas dosing into the feed) (7)  homogenizer (8)  foamed solution flow rate and density meter (9)  pressure nozzle (10)  pressure sensors (P)  temperature sensor (T).

An aqueous solution of the material to be dried with an appropriate amount of surfactant is prepared in the tank (1) with an agitator, and pumped (2) by a thermostated pipeline to the spraying nozzle (10). Foaming gas from a high-

Fig. 6.5 Schematic diagram of a gas-admixing foaming system. See text for details. 6.3 Foam Spray Drying 169 pressure tank (3) is injected into the feed in the mixing device (7). The amount of supplied foaming gas (e.g., nitrogen) must be measured and controlled by means of, for example, a mass flow meter (6). A mechanical homogenizer (8) should be used to produce a uniform mixture of foaming gas and raw material and to eliminate pulsations of flow. From an economic point of view, foaming reduces the cost of manufacturing instant powders by requiring only minor equipment changes that do not affect the efficiency of a dryer to produce standard types of powder. Abdul-Rahman et al. (1971) conducted fundamental experiments on the drying of a single drop of a 15% w/w aqueous solution of sodium caseinate foamed with N2 and NO2. The suspended foamed drop, with an initial mass of either 1 or 2 mg, was exposed to bone-dry air at 100 C and 200 C, respectively, À flowing with a relative velocity of 0.94 m s 1. The results of these experiments showed that the drying and temperature histories of foamed drops were similar to those of non-foamed drops. The foamed drops, which contained approxi- mately 56% of inert gas (volume fraction), exhibited increased drying rates when compared to non-foamed drops. Although foaming did not suppress the temperature level, the reduced exposure time of the drop, which resulted from the accelerated drying rates, indicated that a five- to 10-fold reduction of thermal degradation could be expected for foamed drops. The results of experiments with drops foamed to different volume fractions of nitrogen suggested that the drying time would be shortened with an increased volume fraction of foaming gas. This proposal concurs with a note by Crosby and Weyl (1977), who stated that drying would be promoted when the gas fraction in foam spray-dried feed stocks was relatively high. Subsequent data delivered by Crosby and Weyl (1977) on the moisture change of foamed and non-foamed supported single drops (15% w/w aqueous sodium caseinate solution) showed a quicker dehydration rate for the foamed drops (Fig. 6.6). The effect of foaming on spray parameters and losses of volatile organic components during foam spray drying was investigated by Frey and King (1986). The tests performed indicated that, for the gas-admixing method of foaming, foaming affected the initial atomization parameters due to an earlier break-up of the liquid sheet to droplets (reduction in sheet length) compared to the non-foamed solution. However, this effect was not observed for the gas- desorption method. Frey and King (1986) proved that foaming by gas-admixing enhanced the retention of highly volatile substances (Fig. 6.7) and also improved powder solubility. Additionally, Frey and King (1986) compared the retention of highly volatile substances in gas-admixing and gas-desorption foam spray drying, and concluded that an enhanced retention of volatile substances takes place only in the gas-admixing method, as the result of an earlier break-up of the liquid flow to droplets during atomization. A noticeable decrease in n-propyl acetate retention was found for the gas desorption feed foaming method when compared to non- foamed feed; this was due to the nucleation of bubbles within the droplets, which promoted the escape of volatile compounds (Fig. 6.8). 170 6 Drying of Foamed Materials

Fig. 6.6 Moisturechangeforfoamedandnon-foamedsupporteddrops(15%w/waqueoussodium caseinate solution, drying temperature 100 C). According to Crosby and Weyl (1977).

A systematic analysis of the foam spray drying process, including investigations of spray hydrodynamics (particle size distribution, particle velocity, centricity), heat and mass transfer between the phases, drying kinetics and the effects of feed foaming on final product properties, was carried out by Zbicinski and Rabaeva (2010) and Rabaeva (2012).

Fig. 6.7 Retention of n-propyl acetate in the nozzle region during gas-admixing and gas-desorption foam spray drying of 60% sucrose solution (200 C air temperature, 70 bar atomization pressure). Adapted from Frey and King (1986). 6.3 Foam Spray Drying 171

Fig. 6.8 Retention of n-propyl acetate during non-foamed and gas-desorption foam spray drying of 60% sucrose solution (200 C, 70 bar atomization pressure). Adapted from Frey and King (1986).

Figures 6.9 and 6.10 show changes in the moisture content of foamed and non- foamed feed as a function of the distance from the atomizer for two drying air temperatures (175 and 150 C) for the drying of maltodextrin. The drying process À1 proceeded faster for foamed feed (feed rate 9 kg h , foaming gas (N2) rate À 40 g h 1) at both air temperatures. The foam spray drying process of maltodextrin was complete at about 0.4 m (1.2 m for non-foamed feed) from the nozzle at a

Fig. 6.9 Profile of material moisture content as a function of time for spray drying with and without  gas-admixing (maltodextrin, surfactant GMS, foaming gas N2, drying temperature 175 C). 172 6 Drying of Foamed Materials

Fig. 6.10 Profile of material moisture content as a function of time for spray drying with and without gas-admixing (maltodextrin, surfactant GMS, foaming gas N2,drying temperature 150 C).

drying temperature of 175 C (Fig. 6.9) and at about 1 m (1.8 m for non-foamed feed) from the nozzle at a drying temperature of 150 C (Fig. 6.10). The faster fall in foamed material moisture content was explained by an increase in the evaporation area per unit of drop/particle mass due to the expansion of gas trapped by the liquid (Hanrahan and Webb, 1961; Crosby and Weyl, 1977). No effect of feed foaming was reported on the spray hydrodynamics for all admixed gas flow rates.

6.3.2 Final Product Properties

All available studies revealed a strong effect of feed foaming on the final product properties. As early as the 1960s, Hanrahan and Webb (1961) had obtained a stable non- caking powder of cottage cheese whey by introducing compressed air into the solution before atomization. The powder obtained as a result of foaming had excellent flow characteristics, whereas the spray-dried powder that had not been subjected to foaming showed a tendency to adhere to the dryer surface during the process. Information on the effect of foaming on the final properties of spray-dried cottage cheese whey is provided in Table 6.1. These authors found that powder moisture content decreases as a result of foaming, preventing the lactose crystallization that is important in food manufacture where lactose and lactic acid are desirable ingredients. The main goal of the experiments performed by Hanrahan et al. (1962) was to obtain a highly soluble powder which could be reconstituted to a beverage that was

174 6 Drying of Foamed Materials

Fig. 6.11 Particle size distribution in spray-dried powders of whole milk produced with and without gas incorporation. Adapted from Hanrahan et al. (1962).

virtually indistinguishable from fresh pasteurized milk. These authors showed that inert gas, when introduced into the droplets, caused an increase in powder particle diameter and porosity, which in turn affected the material bulk density. The effect of foaming on the particle size distribution of whole-milk powder is shown in Fig. 6.11. Changes in the final product properties of whole-milk foamed with nitrogen, as presented by Hanrahan et al. (1962), are listed in Tab. 6.2. An analysis of the data in Tab. 6.2 shows a decrease in the bulk density of the product, increases in the free fat content and dispersibility, and also an increase in the average individual particle diameter at a higher nitrogen injection rate. A method of producing readily dispersible non-fat dry skim milk by injecting compressed air into the feed was compared with a conventional spray drying process by Bell et al. (1963). These authors studied the effect of various drying conditions on the properties of foam spray-dried non-fat milk, and suggested that the foam spray drying method would produce a rapidly dispersible product during drying, without a subsequent agglomerating step. Bell et al. (1963) claimed that, since the injection of gas improves the efficiency of water removal and enables the drying of skim milk concentrates with up to 60% solids, the dryer throughput might be increased. Tamsma et al. (1971) failed to observe any significant variations in physical properties or flavor when using different foaming gases in the spray drying of whole-milk, and concluded that air can be substituted for nitrogen as the foaming agent. Extensive experiments on the effects of feed rate and admixed gas flow rate on the final product properties for skin-forming materials (maltodextrin) and porous

176 6 Drying of Foamed Materials

materials (detergent) were carried out in a concurrent spray drying tower by Rabaeva (2012). The final properties of maltodextrin and detergent powder for different drying conditions are presented in Tabs. 6.3–6.5. The analysis of results for maltodextrin showed that there is a clear trend in changes of the properties of this product with increasing foaming gas flow rates (Tabs. 6.3 and 6.4). Whilst the bulk, tap and apparent density was decreased, the Sauter mean diameter and particle porosity were each increased, mainly as a result of the expansion of gas bubbles that had become trapped in the droplets during foaming. The wettability of maltodextrin was determined according to the Polish Standard (PN-88/A-86030) and expressed as the time after which a powder was submerged in the water. It was found that the wettability of maltodextrin decreases with an increase in foaming gas flow rate, as larger particles produced by foaming require longer submergence times. A foaming agent (glycerol monostearate; GMS), which was added to the raw material to make the foam more stable, can also affect the wettability properties of maltodextrin. On analyzing the wettability of the detergent, which is a porous material, an opposite effect can be observed, namely that the wettability increases as the wetting time decreases at a higher foaming gas flow rate (see Tab. 6.5). Typically, a better wettability of a porous material is the result of capillary forces that make the wetting of such particles potentially easier than the wetting of skin-forming products, where the skin on the particle surface must be dissolved first. A different effect of feed foaming was observed on the apparent particle density of maltodextrin and detergent. For maltodextrin, the apparent density decreased with higher foaming gas flow rates due to the presence of closed internal voids inside the particles as a result of foaming. For the detergent (Tab. 6.5), the effect of foaming on apparent particle density was less pronounced, due to the porous structure of the material. Although the detergent apparent particle density did not change significantly, its bulk density decreased with an increase in the foaming gas flow rate, which could be related to morphological changes that had occurred during drying (particle shape, surface structure, etc.) No significant effect of feed foaming on the solubility of maltodextrin was observed in the tests performed, whereas a lower solubility of the detergent could be explained by its chemical composition with poorly soluble components in the formula. The natural angle of repose, which determined the “looseness” of powders, is not significantly affected by feed foaming. A slight reduction in the natural angle of repose for maltodextrin was observed with an increase of the foaming gas flow rate (Tabs. 6.3 and 6.4), which was caused by the formation of more cohesive particles (Teunou et al., 1995). Figures 6.12–6.15 show example electron microscopy images of foamed and non- foamed detergent and maltodextrin powders for different drying temperatures. In these images, a significant difference can be observed in the structure of foamed and non-foamed spray-dried products, with foaming leading to an increase in particle diameter and the formation of a “porous” shell. Under certain drying

6.3 Foam Spray Drying 179

Fig. 6.12 Microscopic image of spray-dried powder (detergent, drying temperature 200 C) (Rabaeva and Zbicinski, 2010). (a) Non-foamed; (b) Foamed.

conditions, some of the particles have bigger holes in the shell (Figs. 6.14b and 6.15b), caused by expansion of the nitrogen. A similar observation was made by Hanrahan et al. (1962) and Hassenkloever and Eggers (2010) when investigating milk and coffee powders that had been foamed with nitrogen. Foamed particles of maltodextrin (Figs. 6.13b–6.15b) have a more uniform structure and spherical shape which affects the bulk and apparent densities. For foamed material, the larger the powder particles the smaller is the bulk density of the product (Tabs. 6.3 and 6.4). For non-foamed material, the voids between large particles in the bed are filled with smaller particles, which causes an increase in the

Fig. 6.13 Microscopic image of spray-dried powder (maltodextrin, surfactant GMS, drying temperature 200 C). (a) Non-foamed; (b) Foamed. 180 6 Drying of Foamed Materials

Fig. 6.14 Microscopic image of spray-dried powder (maltodextrin, surfactant GMS, drying temperature 175 C). (a) Non-foamed; (b) Foamed.

bulk density. An increase in the drying air temperature results in an increase of the volume of foaming gas trapped in the particles, which in turn causes a decrease in the apparent particle density (Greenwald and King, 1981; Verhey, 1973). In the case of a detergent, the particles show more damage due to foaming (Fig. 6.12b), and this may result in a better wettability due to an easier water penetration of the burst particles. In powders obtained at the highest foaming gas flow rate there will be a number of fractured particles that would cause a looser packing of the bed of material and a reduction in the bulk density.

Fig. 6.15 Microscopic image of spray-dried powder (maltodextrin, surfactant GMS, drying temperature 150 C). (a) Non-foamed; (b) Foamed. 6.4 Foam-Mat Drying 181

Extensive experiments were conducted by Lewandowski et al. (2012) to determine the effect of feed foaming on the efficiency of sunflower oil microencapsulation, and on selected properties of the final product. These experiments were performed in a pilot-plant concurrent spray dryer, using the gas-admixing technique for foaming. The microencapsulation efficiency of the foamed feed was shown to decrease by about 6% for an air temperature of 150 C, and by 25% for a temperature of 200 C at the highest foaming gas injection rate in relation to non-foamed feed. Nevertheless, the microencapsula- tion efficiency of sunflower oil remained at a level of 80% to 60%, which was satisfactory for the spray drying technique. The same group also confirmed a 50% reduction in the apparent density of powders at different foaming rates in comparison to the powder produced from a non-foamed emulsion. It was concluded that foaming of the emulsion would enable the control of selected powder properties while retaining a high efficiency of the spray drying microencapsulation process. To summarize, the properties of dry powders, belonging to different groups of materials, depend on the parameters of foaming, atomization, and drying. All results presented to date have confirmed a strong effect of feed foaming on the final product properties of porous and skin-forming materials. Examples of these include an increase in the Sauter mean diameter and porosity of the products, and a decrease in the bulk and tap densities at higher foaming gas flow rates. Hence, feed foaming can be used to control the physical properties of selected products in industrial spray drying applications.

6.4 Foam-Mat Drying

In foam-mat drying, liquid foods are first whipped into a stable foam that is then air-, freeze-, vacuum-, or microwave-dried. In general, wet foams are dried by convection (e.g., belt, drum, tray, oven dryers) by flowing hot air over or through a relatively thin (3–10 mm) layer of the foamed material. A benefit of foam-mat drying is that it allows the processing of difficult-to-dry materials such as tomato paste, fruit pulps and juices at low temperatures and within short drying times when compared to non-foamed materials dried using the same type of dryer. The main drawback of the foam-mat drying process is the limited throughput due to the small load of material that can be applied per unit dryer surface area, though this might be compensated by a shorter drying time. The classical drying of food by convection is a slow process that can result in thermal degradation of the material’s structure, shrinkage, a loss of color, and a deterioration of the organoleptic properties of the materials being dried. In the case of ceramic or building materials an increase in both transverse and longitudinal stresses is observed, which can lead to fractures in the material structure. The use of a foam-mat drying technique reduces the energy consumption of the process when compared to the traditional drying of non-foamed material, and also 182 6 Drying of Foamed Materials

intensifies the mass transfer (Hart et al., 1963). Products dried with the foam-mat drying method are characterized by a porous structure (Kadam et al., 2012), a low density, and a low water activity. Due to its shorter drying times, foam-mat drying can also reduce investment and operational costs (Ratti and Kudra, 2006). In foam-mat drying, solutions of foams are first prepared in tanks to which

foaming gas (N2,NO2,CO2 or air) is applied. Subsequently, depending on the amount of foaming gas applied, the quantity and type of surfactant, and also the type of foamed solution, various foams with different structures will be obtained. Generally, a uniform foam consistency is obtained by using high-speed stirrers and homogenizers. Materials with a low moisture content are hydrated to obtain pulp with a 40–60% water content. The process of foam-mat drying requires the use of foam-forming substances that are capable of producing foams of high stability. For food processing, surfactants of biological character belonging to protein groups, such as egg albumin or milk protein, are used. Typically, a mass fraction of 2–10% egg albumin in the material will cause a 40–50% decrease in material density (Rajkumar et al., 2007; Dattatreya et al., 2010; Krasaekoopt and Bhatia, 2012). Subsequently, Borcherding et al. (2008) showed that the addition of 0.5% methylcellulose to mango pulp produced a foam with between 10% and 25% greater stability. Other materials used to create foams to be dried with foam-mat drying include surfactants such as Spam 60, Tween À 80 and GMS. These foams have low density, in the range of 200 to 900 kg m 3. One other parameter that can influence the final product quality is the drying temperature. Depending on the type of material to be dried, and also its thickness, the range of temperatures used for foam drying is 45 to 90 C (Krasaekoopt and Bhatia, 2012). In the case of tomato juice or foamed mango and apple juice containing 40–50% moisture, the temperature range was 50 to 90 C (Kadam et al., 2012; Karim and Wai, 1999; Dattatreya et al., 2010), whereas for dairy products such as foamed yogurt the optimum range of drying temperatures was 50 to 70 C. An increase in the drying temperature not only intensifies the process of evaporation but also increases the susceptibility of the material to deformation. In the case of materials with lower moisture contents, such as foamed bananas (Thuwapani- chayanan et al., 2008), the optimum temperature was 60 C. In order to improve final product quality, it is important to first select an appropriate thickness of foam solution to be dried. Initially, it was shown that, depending on the foam layer thickness, the density and porosity of the structure, and also its moisture evaporation rate, were changed (Bates, 1964). In studies performed by Rajkumar et al. (2007), Sankat and Castaigne (2004), Kudra and Ratti (2008), the thickness of the material dried by foam-mat drying was found to determine the drying time. Typically, foams of 1 to 5 mm thickness evaporated faster than non-foamed materials; as examples, the drying time of a 79.6% mango solution of density À 1020 kg m 3 was 90–180 min, whereas for a foamed solution of the same initial À moisture content and density 540 kg m 3 the drying time was 40–80 min. In the case of 10–20 mm-thick foamed banana pulp, the drying time was reduced from 46 h to 10 h (Sankat and Castaigne, 2004, Kudra and Ratti, 2008). Due to the high sugar content of these materials, the use of foams thicker than 20 mm may lead to 6.4 Foam-Mat Drying 183 physical changes in the final product. However, in order to eliminate any negative effects of the sugars and acids, which cause product caramelization at higher temperatures, the level of such components should be reduced in the material to be dried (e.g., Hart et al., 1963; Sankat and Castaigne, 2004). The drying of non-foamed fruit juice generally results in the production of a high-viscosity liquid syrup, as the high drying temperature employed causes the sugars contained in the material to melt. Foaming facilitates reductions of both the drying agent temperature and drying time, which may have significant effects not only on the costs of dry material production (Ratti and Kudra, 2006) but also the final product quality. The main goal of research into foam-mat drying is first, to determine optimal drying conditions and, second, to select an appropriate surfactant in order to obtain high-quality products (e.g., Vernon-Cartera et al., 2001; Rajkumar et al., 2007; Kadam et al., 2012; Thuwapanichayanan et al., 2012). In general, foamed materials may dry in constant and falling drying rate period, or in falling rate period only. An example of a foam-mat drying process which is controlled by internal diffusion is shown in Fig. 6.16 (Kudra and Ratti, 2006), where no clear period of constant drying rate is apparent on the drying curves for either foamed or non-foamed apple juice when dried as a 19 mm layer (air temperature: 55 C). Similar results were reported by Thuwapanichayanan et al. (2012) for the drying of banana pulp. When the foam drying process is controlled by both convection and diffusion (e.g., Lewicki, 1975, 2006), the drying rate is seen to depend on product characteristics such as the chemical composition, the type of foaming agent, solids content, and the initial moisture content. Drying curves for foamed tomato paste (Lewicki, 1975, 2006) at different foam densities are shown in Figure 6.17, an

Fig. 6.16 Drying curves for foamed and non-foamed apple juice dried in 19 mm layer by air at 55 C. Adapted from Kudra and Ratti (2006). 184 6 Drying of Foamed Materials

Fig. 6.17 Drying curves of foamed tomato paste with both constant and falling drying rate periods. Air temperature 60 C; layer thickness 3 mm. Adapted from Lewicki (2006).

analysis of which shows that the lower the density of the foamed product, the higher is the drying rate (in the case analyzed, the rate is increased by about 50%). Higher drying rates obtained at a lower density result from the specific structure of low-density foams, which contain very small bubbles but have a high level of homogeneity. Aside from layer thickness, air temperature and air velocity, the drying rate of foamed materials depends mainly on the foam’s characteristics such as stability, density, and bubble size. The foam-mat drying technique was used successfully to dry exotic fruits such as mango (Cooke et al., 1976; Rajkumar et al., 2007), starfruit (Karim and Wai, 1999), banana (Sankat and Castaigne, 2004; Thuwapanichayanan et al., 2008, 2012), tomato paste (Lewicki, 1975, 2006; Kadam et al., 2012), and yogurt (Krasaekoopt and Bhatia, 2012). Because the density of foamed materials is lower than that of À non-foamed materials (typically in the range of 300–600 kg m 3), the mass load of a foam-mat dryer is also lower, which in turn leads to reductions in both dryer throughput and drying time. When Marques et al. (2006) compared different methods for drying tropical fruit pulps, they recommended convective foam-mat drying for liquid foods as well as semi-liquid and thermosensitive foods, and stressed that the main advantages of this method were a reduced drying time, quick rehydration, and low drying temperatures. The main drawbacks of foam-mat drying were nonenzymatic darkening, a loss of aromatic components, and problems with foam stability. Kudra and Ratti (2006) performed an energy and cost analysis of convective drying (belt conveyor and drum dryer) of mats of foamed and non-foamed apple juice. The energy consumption for drying foamed apple juice was estimated to be 6.4 Foam-Mat Drying 185

Fig. 6.18 Drying efficiency over moisture content for foamed and non-foamed apple juice (layer thickness 19 mm; air temperature 55 C). Adapted from Kudra and Ratti (2006).

20% of that for drying non-foamed juice, and with an 18% higher throughput that would reduce capital costs by about 11% for a belt conveyor dryer, and by 10% for a drum dryer. According to these authors, the main factor responsible for the reduced drying time was the increased interfacial area of the foamed materials. The drying efficiency for foamed and non-foamed apple juice dried in a 19 mm layer at an air temperature of 55 C is illustrated in Figure 6.18. The drying efficiency of non-foamed apple juice was shown to be about 30% (a standard value for convective drying), whilst the efficiency of foam-mat drying of apple juice was well above 60%, and the result of a well-developed heat and mass transfer area. Thuwapanichayanan et al. (2012) determined the drying curves of banana foams produced with aid from one of three foaming agents, namely egg albumin (EA), soy protein isolate (SPI) and whey protein concentrate (WPC) (Fig. 6.19). The results of these studies showed a strong effect of surfactant type on drying rate due to the development of different foam structures, while the exponential decrease in moisture content with time proved that the drying process of banana foams was controlled by internal diffusion. Recently, several investigations of foam-mat freeze-drying have been reported to combine the advantages of foam-mat drying and freeze-drying, notably a relatively short drying time and the production of high-quality products. The main areas of application included mainly the pharmaceutical and medical industries (Whang et al., 1995; Lee et al., 1997; Kim et al., 2004; Chino and Dunand, 2008), as well as the food and ceramics industries (Fukasawa et al., 2001; Marques et al., 2006; Muthukumaran et al., 2008; Raharitsifa and Ratti, 2010). Muthukumaran et al. (2008) studied the foam-mat freeze-drying of egg white with xanthan gum as a stabilizer to estimate the total drying time and quality of the 186 6 Drying of Foamed Materials

Fig. 6.19 Drying curves of banana foams with different foaming additives. Egg albumin (EA), soy protein isolate (SPI), and whey protein concentrate (WPC). Adapted from Thuwapanichayanan et al. (2012).

powder (Fig. 6.20). On analyzing the moisture content reduction, it could be seen that almost 15 h was necessary to remove 50% of ice from the non-foamed material, compared to only 6 h with the foamed sample. This was attributed to an increase in the diffusion coefficient due to the development of a porous structure of the foamed sample.

Fig. 6.20 Foam-mat freeze-drying of egg white. Stabilizer, xanthan gum. Adapted from Muthukumaran et al. (2008). 6.5 Summary 187

Raharitsifa and Ratti (2010) showed that foaming reduced the time of the freeze- drying process for the same sample thickness in comparison with drying non- foamed apple juice. The foam-mat freeze-drying technique was also successfully applied to produce nanoporous cellulose foams (Deng et al. (2009) with a porosity of 99%, and also macroporous solid foams from multiwalled carbon nanotubes (Thongprachan et al., 2008). Vacuum foam-mat drying (Jaya and Das, 2004), or a combination of vacuum and freeze-drying (Moy, 1971) or vacuum and microwave drying (Sundaram and Durance, 2008), represent other options to produce high-quality powders with a reduced oxidative degradation (i.e., browning), a high storage stability, and free- flowing properties.

6.5 Summary

Foam drying techniques enable the processing of difficult-to-dry materials (e.g., fruit pulps, juices and dairy products) at low temperatures and with short drying times. As a consequence, the thermal degradation of dry products can be reduced when compared to the convective drying of non-foamed materials, whilst the desired final product properties can be obtained. The use of a foam-mat drying technique reduces the energy consumption of the process in relation to the traditional drying of non-foamed material, and also intensifies mass transfer. Products dried with a foam-mat drying technique are characterized by a porous structure, a low density and low water activity; yet, due to the shorter drying times required, the method can also reduce investment and operational costs. The lower throughput of the foam-mat drying process, due to the smaller load of material per unit surface area of the dryer, might be compensated by a shorter drying time. Nonenzymatic darkening of the product, a loss of aromatic components and problems relating to foam stability are the main drawbacks of foam-mat drying. In foam spray drying, essentially all studies performed have confirmed a strong effect of feed foaming on the properties of both porous and skin- forming material products. These include an increase in Sauter mean diameter and porosity of the products, a decreased bulk, tap and apparent particle density, and an influence on particle morphology. Foam spray drying enhances the retention of highly volatile substances and improves powder solubility (for gas-admixing foaming methods). The higher drying rate observed with foam spray drying can be explained by an increase in the evaporation area per unit of drop/particle mass due to an expansion of foaming gas trapped by the liquid. Feed foaming in spray drying enables selected physical properties of the products to be controlled, under industrial conditions. The foam spray and foam-mat drying technologies represent modern solutions aimed at reducing the environmental impact and economic costs of drying by 188 6 Drying of Foamed Materials

enabling the use of noxious materials at decreased energy consumption during the process.

Additional Notation Used in Chapter 6

Nil

Abbreviations

d.m. dry matter EA egg albumin GMS glycerol monostearate SPI soy protein isolate WB wet basis WPC whey protein concentrate

References

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7 Process-Induced Minimization of Mass Transfer Barriers for Improved Drying Henry J€ager, Katharina Sch€ossler, and Dietrich Knorr

7.1 Introduction

The drying of solid food matrices depends heavily on diffusion and mass transfer through the cells and tissue. The presence of intact cell membranes in the food raw material limits mass transfer processes due to their barrier function. Furthermore, the tissue structure and the network of intercellular air spaces affect mass transfer during drying. The application of an external pulsed electric field (PEF), the induced local structural changes of the cell membrane, and also the increase in permeability due to the appearance of permanent membrane pores were found to positively affect any subsequent drying processes of porous plant materials. PEF provides an alternative to the mechanical, thermal or enzymatic cell disintegration of plant tissue, providing a short-term (milliseconds), low-energy pretreatment (<10 kJ kg 1). High-intensity ultrasound has the ability to alter a material’s characteristics due to interactions between the sound wave and the food product at a structural level. Transmission of the ultrasonic wave takes place by the oscillation of medium particles, and the damping of ultrasound energy in deeper tissue layers forms the basis of the modification of tissue structures, leading to an impact on simulta- neously occurring mass and heat transfer processes during drying. Taking advantage of the specific potentials and opportunities of these new processes for the modification of tissue structure before or during drying, including the understanding and control of the complex process–structure relationships, provides the possibility for a science-based development of an improved drying process. In this chapter, relevant background information will be provided regarding the principles of PEF and ultrasound technology, and its impact on biological tissues. The application of PEF as a pretreatment of raw materials prior to osmotic dehydration, air- and freeze-drying, with the resultant effects on drying character- istics and product quality, will be discussed. In the case of ultrasound, concepts for

Modern Drying Technology Volume 5: Process Intensification, First Edition. Edited by Evangelos Tsotsas and Arun S. Mujumdar. Ó 2014 Wiley-VCH Verlag GmbH & Co. KGaA. Published 2014 by Wiley-VCH Verlag GmbH & Co. KGaA. 192 7 Process-Induced Minimization of Mass Transfer Barriers for Improved Drying

its simultaneous application during the aforementioned drying processes will be presented, including relevant aspects regarding system implementation. Attention will be focused on the drying of plant-based, structured raw materials such as fruits and vegetables in order to illustrate underlying principles for the application of PEF and ultrasound. PEF effects, as a pretreatment for tissue dis- integration, are related to cellular structures and therefore are limited to respective raw materials. In the case of contact ultrasound, permanent cell disintegration effects resulting from pretreatment occur only at the sonicated surface; however, such simultaneous application provides additional effects on tissue structure and related mass and heat transfer processes, examples of which are provided with regards to the drying of plant-based, structured raw materials.

7.2 Structural Characterization of Plant Raw Materials and Impact of PEF and Ultrasound

Plant tissues have a complex structure, and their material properties are spatially dependent and highly inhomogeneous. Plant tissue consists of an outer epidermis, parenchyma cells, and supportive tissue. The cells of the epidermis are tightly packed and may contain a waxy material in order to provide protection against biological or physical stresses. The parenchyma often includes an extensive network of air spaces and consists of semi-rigid polyhedral cells. Differences in the characteristics of the network of air spaces and the resultant differences in levels of porosity will impact on the rates of water and heat transfer during the drying process. Porosity levels for apple tissue fall within the interval of 14–25%, whereas for many other fresh plant materials the porosity values are lower; for example, 4% for carrot and 2% for potato (Karathanos et al., 1996). This in turn will affect the diffusion coefficient, which is 10 9 m2 s 1 for more porous apple tissue, compared to 10 11 m2 s 1 for potato and carrot (Okos et al., 1992). It may also affect the impact of PEF and ultrasound, as will be discussed in following sections. The vacuole comprising most mature plant cells contains secondary metabo- lites and enzymes. The cell envelope is composed of cellulosic cell walls bound by pectinaceous middle lamella. A phospholipid-bilayer membrane is located on the inside of the cell wall (plasmalemma), as well as around the vacuole (tonoplast). Disruption of the cell or vacuole membrane leads to the release of hydrolytic enzymes which cause the degradation of pectins or cellulose. This in turn weakens the cell wall and tissue structure and leads to a loss of firmness and crispiness due to a dissipation of the intracellular osmotic pressure (turgor) (Taiz and Zeiger, 2007).

7.2.1 Methods for Analysis of Tissue Structure and Quantification of Cell Damage

Several methods have been proposed to study and to quantify the degree of membrane permeabilization and tissue modification. The damage degree is 7.2 Structural Characterization of Plant Raw Materials and Impact of PEF and Ultrasound 193 defined as the ratio of damaged cells to the total number of cells. A microscopic analysis of PEF-treated tissue stained with neutral red was applied successfully by Fincan and Dejmek (2002). By regarding plant tissue as a semi-solid, multiphase, heterogeneous material, these authors emphasized the need to characterize the desired level of cell permeabilization in order to optimize PEF treatments, and to elucidate effects such as mass transport and mixing of the vacuole content, the cytoplasm and the liquid of the apoplastic space, as well as their interaction with the cell wall. More complex staining methods using selective and fluorescent dyes have been used by Phoon et al. (2008), in combination with microscopic analysis. The measurement of cell membrane permeabilization of single cells or protoplasts, using fluorescent dyes, is also possible by means of flow cytometry (Shapiro, 2003), and this technique is gaining increasing interest for the characterization of plant cell microstructures. As membrane permeabilization affects the turgor of the cell, measurements of the latter (Tomos, 2000) and of texture characteristics such as stress deformation, as well as relaxation tests of complex tissues, may be employed in order to evaluate the degree of cell rupture and tissue damage (Lebovka et al., 2004). Other simple methods for analyzing the degree of cell rupture might entail measurements of the release of plant pigments from permeabilized cells (Dornenburg€ and Knorr, 1993), or the release of ions into the cell-containing media (Saulis et al., 2007). Knorr and Angersbach (1998) quantified the liquid release of permeabilized cells by subjecting the tissue to a centrifugal force that causes liquid release to occur from PEF-treated cells, but not from undamaged cells. The ability to estimate the degree of damage by using diffusion coefficient measurements in PEF-treated biological tissue was suggested by Jemai and Vorobiev (2002). The degree of damage can be determined from solute extraction or convective drying experiments but diffusion techniques are indirect and invasive for biological objects and may affect the structure of the tissue. By using the electrical conductivity response of plant tissue, cell disintegration criteria can be defined based on impedance measurements (Knorr and Angers- bach, 1998; Lebovka et al., 2001). This method is based on the interfacial polarization effects of the Maxwell–Wagner type at intact membrane interfaces. Measuring conductivity–frequency spectra (in the frequency range from 103 to 107 Hz) allows the degree of membrane disintegration to be quantified. Conductiv- ity changes in the low-frequency ranges are the result of irreversible membrane permeabilization; however, in the high-frequency ranges the conductivity value of intact cells and of cells with ruptured membranes is the same, since in this frequency range the cell membrane does not present any resistance to the electrical current measured (see Fig. 7.1). A successful application of an acoustically derived determination of the degree of cell permeabilization was made by Grimi et al. (2010), who subsequently established a relationship between the acoustic disintegration index and the con- ductivity disintegration index. This nondestructive acoustic technique is especially applicable if fruits or vegetables are processed as whole unpeeled samples, based 194 7 Process-Induced Minimization of Mass Transfer Barriers for Improved Drying

Fig. 7.1 Typical frequency–conductivity spectra of plant tissue with intact cells, partial ruptured cells and totally ruptured cells in the frequency range of the measured current of 1 kHz to 50 MHz. Reproduced with permission from Knorr and Angersbach (1998).

on the modulus of elasticity which is proportional to the index of firmness of the plant tissue. The application of a series of gentle taps to the sample cause it to resonate at a frequency that depends on its weight and firmness. The vibration of the sample is then captured by acoustic measurements and an acoustic spectrum derived which shows the amplitude–frequency correlation for the untreated or PEF-treated sample. An acoustic disintegration index can also be calculated by comparing the pulse response-derived index of firmness of the samples. Cell disintegration may also be achieved using techniques other than PEF or ultrasound, with different mechanisms of action:

Thermal treatment causes denaturation of the membrane proteins as well as an expansion of the intracellular liquid, and this results in a permeabilization of the cell. Cell disintegration as a result of freeze–thawing is based on destruction of the cell envelope by the formation of large ice crystals when applying slow freezing rates. Mechanical cell disintegration by crushing or milling leads to a particle size reduction which in turn results in an increasing degree of cell disintegration as the particle size becomes close to or is less than the cell size.

As membrane structures are also affected by these physical means of cell disintegration, the aforementioned analytical techniques are appropriate. In addition to the characterization of cell structure modification, nondestruc- tive measurements of water content, of local tissue structures and porosity, and of temperatures before and during drying, are required in order to 7.2 Structural Characterization of Plant Raw Materials and Impact of PEF and Ultrasound 195 describe any peculiarities during the drying of PEF-pretreated material or during ultrasound-assisted drying. Magnetic resonance imaging (MRI) was introduced as a means of monitoring the drying of food, and the information obtained can be related to water distribution and porosity, based on the proton concentration that depends on the water concentration in the tissue (Hardy, 2012). Thermally induced chemical shifts of the proton resonance frequency can be used for the measurement of local temperatures (Hindman, 1966; Knoerzer et al., 2009). Noninvasive measurements, combined with the three- dimensional resolution of water and temperature distribution and tissue porosity, provide relevant information that could be applied to quantify related properties and effects (Regier et al., 2007).

7.2.2 PEF: Principles and Impact on Plant Tissue Structure

7.2.2.1 Introduction to PEF Technology PEF technology can be used to induce the non-thermal permeabilization of cell membranes and, depending on the treatment intensity (external electric field strength, number and duration of the electric pulses) and cell properties (size, shape, orientation, conductivity), the pore formation may be either permanent or temporary (Zimmermann et al., 1974, 1976; Angersbach et al., 2000). The treatment involves the application of very short electric pulses (1–100 ms) at electric field intensities in the range of 0.1–1kVcm 1 for the reversible per- meabilization by means of stress induction in plant cells, 0.5–3kVcm 1 for the irreversible permeabilization of plant and animal tissues, and 15–40 kVcm 1 for the irreversible permeabilization of microbial cells. The aforementioned field intensities lead to the formation of a critical transmembrane potential, which is regarded as being the precondition for cell membrane breakdown and electropora- tion (Tsong, 1996). Irreversible electroporation results in a loss of turgor, a leakage of cytoplasmic content, and lysis (Rubinsky, 2010), whereas reversible permeabilization leads to the formation of conductive channels across the cell membrane, though the electrically insulating properties will recover within seconds (Glaser et al., 1988; Angersbach et al., 2000). A typical PEF system consists of a pulse modulator which supplies high- voltage pulses of sufficient peak voltage and repetition rate, and a treatment chamber in which the product is exposed to the electric field. For pumpable media such as fruit juices, milk or fruit mashes, the treatment can be performed continuously in a tubular treatment chamber consisting of cylind- rical electrodes. Whole fruits or pieces can be submerged in water in a flume, but a dry PEF treatment with direct contact between the product and the electrodes is desirable for meat pieces, in order to prevent cross-contamination. Further information on the principles and mechanisms of action of PEF, as well as its technical aspects and applications, are available elsewhere (Lelieveld et al., 2007; Raso and Heinz, 2007). 196 7 Process-Induced Minimization of Mass Transfer Barriers for Improved Drying

PEF-based cell disintegration is generally applied as separate pretreatment prior to drying, as an additional step of raw material processing. The electric energy input and field strength are selected depending on the type of raw material and the desired level of cell permeabilization. The various processing intensity require- ments determine the treatment costs; for example, field strength will have a major influence on investment costs as it depends on high-voltage power supply equipment, while energy input will primarily determine the cost of the operation as it requires a certain electrical energy consumption. However, even if PEF requires an additional electrical energy input, this may have a beneficial effect on the total energy consumption of mass transfer processes such as extraction, pressing, or drying. As a result, the processing times are reduced, production capacities are improved, and the consumption of both water and raw material is decreased (Toepfl et al., 2006). Taken together, the pulse modulator typology and components, the supplier and the processing and product parameters lead to a wide range of investment costs, with energy requirements for cell disintegration ranging from 1 to 3kWht 1. Industrial-scale equipment with capacities of up to 50 t h 1 for the treatment of plant raw material have recently been developed, based on the installation of a single 80 kW semiconductor-based Marx Generator with a peak voltage of 30–60 kV and repetition rates of up to 300 Hz (S. Toepfl,personal communication). Such a pulse generator would be linked to a treatment chamber where the product is exposed to the electric field. For this purpose, belt systems have been designed and implemented successfully for industrial production systems. The plant raw material (e.g., potato) is submersed in water and transported via a moving conveyor belt through the electric field zone (see Fig. 7.2). As cell disintegration by PEF takes place throughout the tissue, the treatment can be applied to the whole fruit or vegetable as well as to smaller pieces after cutting, depending on the process objectives and the steps and line set-up of the conventional process.

7.2.2.2 PEF: Impact on Plant Tissue Structure The primary objective of applying PEF treatments to solid raw materials is to effect cell disintegration and modify the tissue microstructure and texture, thereby improving and/or minimizing any subsequent processing steps such as drying or extraction (Grimi et al., 2007; Lebovka et al., 2007). In this section, the aim is to discuss the influence of PEF on food, the most relevant effects being a direct impact on food microstructure and textural proper- ties due to the membrane permeabilization that occurs when biological tissues are exposed to an external electric field, together with the related release of cell contents and loss of turgor. On the other hand, the food matrix – namely the plant tissue structure – may affect the PEF performance not only in terms of the achievable level of cell disintegration but also of the impact of cell disintegration on subsequent processing steps. According to Zimmermann et al. (1974), the transmembrane potential is 7.2 Structural Characterization of Plant Raw Materials and Impact of PEF and Ultrasound 197

Fig. 7.2 Belt system for the PEF treatment of (C) the electrical treatment zone with plant raw materials such as whole potato with electrodes. The width of the belt is 800 mm, a capacity of 50 t h1, consisting of (A) a basin and electrode distance is 170 mm. Illustration filled with water, (B) the conveyor belt, and courtesy of S. Toepfl.

3 DV ¼ Ercos q; ð7:1Þ m 2 where cos q describes the orientation of the cell in the electric field of the electric field strength E, and r is the cell radius. Consequently, the transmembrane potential is proportional to the cell size, so that a lower electric field strength will be required for the electroporation of larger cells. Not only cell size but also tissue structure may affect the PEF performance, as the efficacy of cell damage is determined by the local electric field near the cell. The distribution of the local electric field strength inside the inhomogeneous material is a complex function of electrical properties, tissue porosity, and structure (Sahimi, 1994). Differences in porosity exist between various types of raw material, or even within the same food material, which may then cause a gradient of electrical conductivity inside the sample and lead to differences in permeabilization effectiveness. As PEF affects the cell membranes, it can also be expected to influence the texture of products in which the structure is largely dependent on cell integrity. The release of water into the extracellularspace,changesinwaterretention properties, damage to the cell membranes and loss of turgor, as well as a possible enhancement of enzyme reactions, may occur as either immediate or delayed effects. The destruction of cell membranes and removal of cellular turgor were considered to be the main impact factors for the modification of texture and 198 7 Process-Induced Minimization of Mass Transfer Barriers for Improved Drying

viscoelastic properties of plant tissue after PEF treatment (Lebovka et al., 2004). A comparison between thermal and PEF treatment, as well as a combination of both, with regards to the textural properties of carrots, potatoes and apples was conducted. In this case, the application of PEF treatment had only limited effects on cell wall softening and reduction of cellular integrity when compared to thermal treatment. After PEF treatment, the tissues partly lost their initial strength, and both the elasticity modulus and fracture stress were decreased with increasing times of PEF treatment. This effect was most pronounced for carrot and apple tissues, with even highly PEF-disintegrated potato tissue showing almost the same stress–deformation behavior as an intact tissue. Heat treatment influenced not only the turgor but also other characteristics of the tissue, including the inner chemical structure of the cell wall, cell wall breakdown or swelling, starch gelatinization, protein insolubili- zation, or the expulsion of trapped air. The influence of cell permeabilization on the textural properties of plant tissues and the correlation between failure stress and degree of pore formation was studied by Bazhal et al. (2003). Compression tests of electroporated apple discs were performed, and the porosity, pore size distribution, bulk density and skeletal density of the apple tissues after hot-air drying were analyzed. Stress– strain curves for control and PEF-treated samples exhibited a characteristic elastic behavior at low strain levels, and a nonlinear strain response leading to a bioyield failure point above strain levels of about 0.15. This bioyield point was described by the authors as the point at which external pressure and cell turgor pressure are in equilibrium and cell rupture begins. Exposure of the apple tissue to PEF affected the bioyield point, while increasing the number of applied pulses reduced the bioyield stress values. A viscous behavior of the tissue was observed from nonlinear stress–strain dependency at large strain. A sharp decrease in the compression modulus, with an increase in disintegration index, was noted for low strain values due to the electric field-induced pore formation. Even at a low degree of tissue destruction, apple tissues became more com- pressible after PEF treatment. Modifications of the tissue physical properties by PEF, as described above, mainly result from changes in the cell membrane’s functional properties. Investigations conducted by Pereira et al. (2009) revealed the impact of size and persistence of PEF-created membrane pores on the viscoelastic properties of potato tissue, as indicated by values of tan d (the ratio between elastic modulus and viscous modulus) obtained from dynamic rheological measure- ments. More intensive PEF treatment, which resulted in the formation of larger pores and a faster escape of cell liquid, led to a higher increase in tan d. Reversible permeabilization and pore resealing can be observed by an initial slow decrease followed by a rapid increase in the tan d value due to the recovery process and the reabsorbance of leaked cell contents, which in turn increases the turgor pressure. The viscoelastic response was also shown to be independent of the degree of permeabilization, but was affected by the size and distribution of the membrane pores. 7.2 Structural Characterization of Plant Raw Materials and Impact of PEF and Ultrasound 199

7.2.3 Ultrasound: Principles and Impact on Plant Tissue Structure

7.2.3.1 Introduction to Ultrasound Technology The term ultrasound describes sound waves with frequencies above human hearing, in the range from 16 kHz up to several MHz. Whilst sound waves are pressure waves transmitted by an oscillation of medium particles, the generation of ultrasound is based on transducers transforming electrical energy into high- frequency mechanical vibrations. In most industrial systems a generator is used to convert mains electricity into a high-frequency alternating current, which is then converted into high-frequency ultrasonic vibrations by a piezoelectric material. Different emitters can be used in order to transmit these vibrations – the ultrasonic wave – into a medium or product. Ultrasound is used in a variety of industrial applications, with a very wide range of objectives. Low-intensity ultrasound (with typically I < 1Wcm 2)is used for the process control and nondestructive testing of products (Knorr et al., 2004; Leemans and Destain, 2009), whereas high-intensity ultrasound (with I of 10–1000 Wcm 2) is applied to intensify processes, such as the inactivation of microorganisms (Lee et al., 2009), homogenization (Wu et al., 2001a), or drying (Mulet et al., 2003; see also Chapter 8 of this volume). These very different applications underline the fact that the parameters of ultrasound and its interaction with a product must be well understood in order to allow successful process development. Ultrasonic waves can be transmitted through gaseous, liquid, or solid media. The speed of the ultrasound wave, its transmission, reflection and damping, and also its energy, are greatly affected by the characteristics of these transmission media, as well as by objects in the sound field. When a sound wave propagates through a medium it is exponentially dampened depending on the damping coefficient a characteristic of that medium, according to the relationship:

ax A ¼ A0 e ; ð7:2Þ where A is the amplitude at a certain travelled distance x, and A0 is the amplitude of the incident wave (Humphrey, 2007). When passing from one medium to another, sound waves are partly reflected depending on the acoustic impedances of the two media. The acoustic impedance is the relation of sound pressure to particle velocity in a certain medium. The larger the difference between the acoustic impedances of two media, the larger is the reflected part of the sound wave. This reflected part can be expressed by the reflection coefficient (R):  Z Z 2 A R ¼ 1 2 ¼ r ; ð7:3Þ Z1 þ Z2 A0 where Z1 and Z2 are the acoustic impedances of media 1 and 2, respectively, and A0 fl and Ar are the amplitudes of the incident and the re ected sound wave. In contrast, a travelling sound wave itself causes effects in the transmission medium. In liquids, vapor and gas filled microbubbles are affected by the pressure 200 7 Process-Induced Minimization of Mass Transfer Barriers for Improved Drying

variations, which can lead to their oscillation, growth, and collapse (Kentish and Ashokkumar, 2011); this process, termed cavitation, is important for a large variety of industrial ultrasound applications. Shear effects and microstirring occur in the vicinity of the oscillating bubbles, whereas the sudden compression of gas or vapor during their collapse leads to shockwaves and extremely high local temperature and pressure peaks in the range of 104 K and 106 MPa, respectively (Noltingk and Neppiras, 1950). Sonoluminescence and radical formation can be caused by these harsh conditions (Suslick, 2003), while part of the acoustic energy is also absorbed by the sonicated medium and converted into thermal energy. As a consequence, the medium temperature is increased, depending on its absorption coefficient and its specific heat capacity (Humphrey, 2007).

7.2.3.2 Ultrasound: Impact on Plant Tissue Structure The effects of ultrasound on cells range from the stimulation of living cells (e.g., to increase the production of secondary metabolites) to cell disruption, depending on the intensity of the treatment. Cell disruption, which is best studied in relation to the antimicrobial action of ultrasound, occurs due to cavitation and its associated effects (Kinsloe et al., 1954; Hughes and Nyborg, 1962). Synergistic effects have been identified for the combined application of ultrasound and heat or pressure; these are most likely caused by a heat- or pressure-related weakening of the cell membrane, leading to increased sensitivity against cavitation (Patist and Bates, 2008). In addition to microbial inactivation, the cell-disrupting effects of ultrasound have been used successfully for the treatment of plant cells allowing, for example, an increased extraction of secondary metabolites (Wu et al., 2001b). Figure 7.3 shows the impact of ultrasound on cell integrity, by monitoring changes in pH and conductivity in a medium (water) following the ultrasound treatment of cultured apple cells (Malus domestica). The leakage of organic acids and ions provides a measure of the damage to the cell membrane. Leakage, and thus cell disruption, can be seen to increase with the increasing amplitude of the applied ultrasound waves, despite the energy input being constant at 4.7 kJ kg 1. With regards to cavitation effects, higher amplitudes lead to larger pressure differences acting on the medium and plant cells, which in turn results in stronger mechanical forces during the bubble collapse. Consequently, for individual cells treated in liquids, ultrasound can lead to improved drying characteristics due to the lack of a mass transfer barrier in the form of an intact cell membrane. Intact plant tissues can be affected quite differently when treated with ultrasound, as demonstrated when ultrasound was applied (in water) to cylinders cut from fresh apple parenchyma. Although the specificenergyinput was as high as 100 kJ kg 1, this did not impact on either the cell integrity or the drying kinetics (Angersbach et al., 1997) (Fig. 7.4). The cavitation effects in the surrounding water were limited to the surface of the apple cylinder. The physical characteristics of the intact tissue differed from those of water, and caused a significant damping of the sound wave that, ultimately, was responsible for the 7.2 Structural Characterization of Plant Raw Materials and Impact of PEF and Ultrasound 201

Fig. 7.3 Influence of amplitude of vibration bath (Sonorex Digital 108; Bandelin, at the sonotrode tip and treatment time on Germany; f ¼ 35 kHz); all other ultrasound cell leakage of an apple cell suspension treatments: ultrasonic probe system culture at constant energy input of (UIP1000; Hielscher Ultrasonics, Germany; 4.7 kJ kg1.Amplitude3mm: ultrasonic f ¼ 20 kHz).

Fig. 7.4 (a) Drying curves and (b) an ultrasonic probe system (UIP2000; impedance measurement of apple cylinders Hielscher Ultrasonics, Germany; f ¼ 20 kHz). (w 2 cm, height 1 cm), cultivar Braeburn, Drying was carried out at 60 C. Impedance after sonication in 15 ml of water (US) measurement is based on the change in compared to nonsonicated control samples frequency-dependent conductivity of plant (non-US). Sonication was carried out with tissueindependenceofcell an amplitude of 160 mm for 6 s, resulting in permeabilization. a specific energy input of 100 kJ kg1 with 202 7 Process-Induced Minimization of Mass Transfer Barriers for Improved Drying

reduced cell-destructive effect of ultrasound in intact tissue. Such damping also limited the applicability of ultrasound as a pretreatment when compared to the application of PEFs. The drying-enhancing effects of ultrasound have also been demonstrated for the hot-air drying (Gallego-Juarez et al., 1999; García-Perez et al., 2006, 2011; Schossler€ et al., 2012a) and freeze-drying (Bantle et al., 2010; Schossler€ et al., 2012b) of fruits and vegetables, when ultrasound was applied during the drying process (see also Chapter 8). In a study designed to monitor the modification of cell structure and mass transfer in potato tissue during ultrasound application, an ultrasonically enhanced water removal was shown to have occurred even in tissue areas that had been unaffected by ultrasonic cell disruption (Schossler€ et al., 2012c), confirming that other ultrasound-mediated effects had accelerated the water removal. Cells in plant tissue are connected to each other via the middle lamella, which consists of pectic compounds and protein. Plant tissues can vary in their porosity; that is, in the volume fraction of free intercellular space, as the cells in the tissue can be either more densely or less densely packed. Typical examples of the large differences that can occur in raw materials include potato tissue, which is practically nonporous with only about 2% intercellular space, and highly porous apple tissue with up to 20% intercellular space (Lu and Puri, 1991). This porosity has an important effect on the susceptibility of a tissue towards the ultrasound effects (García-Perez et al., 2007, 2011; Schossler€ et al., 2012c). On travelling through a plant tissue, the transmission medium for the sound wave will be changed with each transition, from surrounding medium to cell membrane, from cell membrane to cell plasma, and again to the cell membrane when the sound wave passes the cell. It can be assumed that, in highly porous products such as apple tissue, the sound wave can travel through the pore network and is less damped when arriving at the deeper tissue layers. This effect was visualized by analyzing micrographs of apple and potato tissues in different tissue layers after treatment with an ultrasound probe, using water as the sound-transmitting medium (Schossler€ et al., 2011). Although both tissues were affected at the surface, where strong cell disruption occurred, the dense potato tissue damped the sound wave more efficiently at only 800 mm from the surface, such that no more tissue changes were observed. In contrast, the apple tissue showed slight irregularities in cell shape at this distance, even when sonicated for only half of the time and at a lower amplitude than the potato tissue (Fig. 7.5). However, textural changes were detectable for both products and also in the tissue layers, where no further cell damage was observed on micrograph analysis. Despite an exponential damping of the sound wave travelling through the tissues, the sound energy in tissue layers up to 5 mm from the sonicated surface (the maximum distance analyzed in that study) was sufficient to cause microstructural effects considered responsible for loosening cell cohesion (Schossler€ et al., 2011). These findings suggest that ultrasound- mediated effects would cause an improved water removal during directly assisted drying processes. 7.2 Structural Characterization of Plant Raw Materials and Impact of PEF and Ultrasound 203

Fig. 7.5 Micrographs of (a) apple and (b) probe system (UIP2000; Hielscher potato tissue, untreated (non-US) and Ultrasonics, Germany; f ¼ 20 kHz): 108 mm sonicated in water using a probe system amplitude, 10 s, 630 J cm2 (apple) and (US). For sonicated samples, micrographs 116 mmamplitude,20s,1400Jcm2 taken from the surface and at 800 mm tissue (potato). The samples were stained with depth are compared. Parameters of the astrablue prior to microscopy. Based on ultrasound treatment with an ultrasonic Schossler€ et al. (2011).

In contrast to PEF, ultrasound acts primarily on water removal at the tissue level, but not at the cell level. Only the outer cells can be directly affected by ultrasonic cell disruption (Schossler€ et al., 2012c), where the sound pressure amplitude and ultrasonic heating effects are most important. In deeper cell layers, the amplitude of the pressure wave and its energy are strongly dampened, as shown in Fig. 7.6. However, cell damage occurs in every drying process as a consequence of water removal from the cell membrane. Hence, water removal from damaged cells can be improved by ultrasound waves in several ways (as summarized in Fig. 7.6), including microstructural effects such as a loosening of cell-to-cell bonding (as discussed above). Water removal from cells with damaged cell membranes can also be increased by micro- streaming, improved evaporation in low-pressure cycles (Mulet et al., 2003), and ultrasonic heating (Humphrey, 2007). In the pore network, the ultrasound waves can reduce the boundary layer thickness and improve the removal of moisture from solid–liquid interfaces. Moreover, compressions and rarefac- tions caused by the sound wave, which have been described as the “sponge effect,” may contribute to an improved moisture movement towards the outer layers of the tissue (Fuente-Blanco et al., 2006). As water is entrapped in the intact cells of a plant tissue, ultrasound-mediated effects depend on cell damage due to enhanced dehydration. It is likely, that a cell disintegration by PEF prior to ultrasound-assisted drying would make the 204 7 Process-Induced Minimization of Mass Transfer Barriers for Improved Drying

Fig. 7.6 Contact ultrasound effects on mass bondings;2,improvedwaterremovalfrom transfer at cellular and tissue levels. Effects cells due to microstreaming, improved are shown in relation to the distance from evaporation in the low-pressure cycle and the sonicated surface. Temperature effects ultrasonic heating; 3, reduction of boundary and sound pressure amplitude are damped layer thickness and removal of moisture in deeper tissue layers, as depicted in the from solid/liquid interfaces; 4, compression graph on the left. M, cell membrane; W, cell and rarefaction of the tissue due to the wall;L,middlelamella;I,intercellular sponge effect; 5, accelerated moisture spaces;P,porenetwork.Dottedlines transport through the pore network; 6, cell represent membrane damage due to disruption close to the sound-transmitting dehydration. 1, Loosening of cell-to-cell surface.

ultrasound effect on water removal even more pronounced. Future studies should be focused on such combined treatments in order to achieve maximum benefits for the drying of cell tissue-based products.

7.3 Pulsed Electric Field (PEF) Application as a Pretreatment

The removal of water is one of the oldest-known methods used for food preservation, and today the food industry mostly employs hot-air drying or freeze- drying to produce dehydrated foods. As the drying rate of the product is determined by the rate at which moisture is transferred from the core to the 7.3 Pulsed Electric Field (PEF) Application as a Pretreatment 205 surface, an acceleration of mass transfer would be advantageous with regards to shortening the drying times, reducing energy consumption, improving product quality, and allowing the drying of larger pieces of food. As electropermeabilization of the cell membrane leads to a drastic increase in mass transfer rates, high-intensity PEF pretreatments can be exploited within the drying processes.

7.3.1 Osmotic Dehydration

Osmotic dehydration results in a partial removal of water from food by its immersion in a hypertonic solution containing sugar or salt. This involves a countercurrent mass transfer process, in which water is drained from the interior to the hypertonic solution, while the solute flows into the food. Osmotic dehydration can be used either alone or as a pretreatment before thermal drying. As the process functions at ambient temperature with the osmotic pressure difference serving as the driving force for water removal, heat damage to the tissue can be minimized. Generally, osmotic dehydration times will depend, among other factors, on the permeability of the cell membranes and the tissue structure. Electroporation of the cell membrane was found to have a positive effect on this mass transfer process. When water-loss rates and the respective diffusion coefficients of carrots during osmotic dehydration (50 Brix sucrose solution at 40 C) were investigated by Rastogi et al. (1999), pretreatment of the carrots with an electric field intensity in the range of 0.22–1.60 kVcm 1 and a total specific energy input of 0.04–2.25 kJ kg 1 resulted in an increased diffusion coefficient and a faster water loss. The maximum treatment intensity caused a cell disintegration index of 0.84, while the effective diffusion coefficients of water and solute, determined using a Fickian diffusion model, were increased exponentially with increasing electric field strength in the above-mentioned range. The level of cell disintegration related to the fraction of ruptured area of the cell membrane, with a rupture of approximately 0.008% of the membrane of cellular materials being sufficient to positively enhance transfer rates during osmotic dehydration under atmospheric pressure conditions. This was achieved with an electric field treatment at 1 kVcm 1. A combined PEF treatment and osmotic dehydration of red bell pepper in a solution of sucrose and sodium chloride was applied as a pretreatment prior to hot- air drying (fluidized bed, 60 C, air velocity 1 m s 1), and had significant effects on the mass transfer rates (Ade-Omowaye et al., 2003a). Effective water diffusivity values for PEF-pretreated samples that had been preconcentrated in osmotic solution for 30 or 60 min ranged from 0.94 to 1.36 10 9 m2 s 1, depending on the treatment conditions, whereas untreated samples showed values from 0.87 to 0.99 10 9 m2 s 1. Red bell pepper, when subjected to a PEF pretreatment and subsequent osmotic dehydration, showed an enhanced rate of mass transfer as indicated by an increase of 11–25% in water loss and of 2–5% in solids gain, when compared to untreated 206 7 Process-Induced Minimization of Mass Transfer Barriers for Improved Drying

samples (Ade-Omowaye et al., 2003b); moreover, the PEF-mediated effect was found to differ between the plasma membrane and the tonoplast (Angersbach et al., 1999). Selective membrane permeabilization may result in structural alterations that favor the transfer of water across the membrane, while limiting solid gain. A minimal solute gain was observed by Heng et al. (1990) when the water outflow was rapid; these authors suggested that effects which favor dewatering may enhance the formation of a concentrated superficial layer that minimizes solute uptake. Subsequent investigations conducted by Lazarides and Mavroudis (1996) revealed that this layer of solute, derived from the soaking solution, disturbs the osmotic pressure gradient and this leads to a decrease in the apparent diffusivity of the solute. The modification of cell structure by PEF, in combination with osmotic dehydration, may have specific advantages for the production of intermediate- moisture foods with minimal solute uptake, though further clarification of the mechanisms involved is required. The osmotic dehydration of PEF-pretreated apple tissue was evaluated with regards to water loss and solids gain by Amami et al. (2005). Both parameters were increased in value, with an increase in electric field strength and pulse number up to an optimal level of 0.9 kVcm 1 and 750 pulses of 100 ms with a specific energy input of 13.5 kJ kg 1. When osmotic dehydration was performed at 25 Cina 44.5% sucrose solution, the kinetics of water and solute transfer showed that PEF pretreatment had increased the diffusion rates. This finding was later confirmed, whereby a Fick’s unsteady-state diffusion equation was used additionally to model mass transfer during the osmotic dehydration of PEF-pretreated apple tissue (Amami et al., 2006). Effective diffusion coefficients of water and solute were calculated and showed an increase due to PEF pretreatment. However, as revealed previously by Ade-Omowaye et al. (2003b), the effect of PEF was more pronounced for the water loss than for the solids gain. The application of PEF as a pretreatment to centrifugal osmotic dehydration showed beneficial effects on water removal, but increased also the solids gain (Amami et al., 2007). When the solute concentration and temperature, as well as centrifugal acceleration, were studied in combination with PEF treatment, synergistic effects were revealed. In further investigations by Amami et al. (2008), PEF and centrifugal osmotic pretreatment were employed as preprocessing steps prior to the convective drying of carrot tissue. Notably, PEF showed a potential to reduce the total time of dehydration operations, since effects on both osmotic dehydration and hot-air drying were observed. PEF pretreatment, in combination with osmotic dehydration, caused the moisture effective diffusivity to be increased from 1.17 10 9 m2 s 1 to 5.1 10 9 m2 s 1, whereas osmotic dehydration alone resulted in a diffusivity of 3.85 10 9 m2 s 1.

7.3.2 Air Drying

Moisture transfer during hot-air drying occurs in two successive steps: (i) from the interior of the tissue to the surface (internal moisture transfer); and (ii) 7.3 Pulsed Electric Field (PEF) Application as a Pretreatment 207 from the surface to the gas bulk (surface moisture transfer). In the second drying stage, the internal moisture transfer is lower than the surface moisture transfer, and a drying front is moving towards the center of the product. In this stage, the cellular and tissue structure plays a major role, since it represents a heat and mass transfer barrier. Heat and mass transfer depend not only on process parameters but also on the raw material characteristics (Gehrmann et al., 2009). Large differences in local diffusion coefficients occur during the removal of intracellular, extracellular and capillary water affected by cellular structure, capillary size and the concentration of dissolved solids. Hence, there is a nonuniform distribution of temperature and moisture, and limitations of mass and/or heat transfer will result from food structures as well as the drying process parameters (Kessler, 2002). PEF-enhanced permeability, which resulted in an improved mass transfer during air drying, and therefore a shorter drying time, was demonstrated for potato tissue by Lebovka et al. (2007). These authors found that the effective diffusion coefficient was a function of the cell disintegration index. Subse- quently, freeze-thawing, PEF pretreatment and thermal pretreatment were each compared with regards to their impact on the drying process. Typically, freeze- thawed samples showed the highest values of diffusion coefficient, whilst thermal pretreatment caused beneficial effects only at mild temperatures (50 C) rather than at the high temperatures (70 C) that might be associated with starch gelatinization. Both, thermal pretreatment at 50 CandPEFpretreatment were shown to result in a similar increase in diffusion coefficient during hot-air drying. The same authors also applied PEF treatment during drying after a predrying period of 1000 s in order to demonstrate the immediate impact of such treatment on the kinetics of the process. The first attempts to study moisture distribution in PEF pretreated tissues during drying were undertaken by Knorr and Angersbach (1998). In this case, a comparison of moisture distributions within potato tissue during fluidized-bed drying with an air flow rate of 2 m s 1 and a drying temperature of 70 C revealed a faster movement of the drying front towards the center of the tissue, and also showed a more homogeneous moisture distribution in the PEF pretreated samples. In convective hot-air drying of red bell pepper at 60 C with an airflow of 1 m s 1, the drying time was reduced by 20–30% for PEF pretreated samples (Ade- Omowaye et al., 2001). Hot-air drying at 70 C of carrot pretreated by PEF (1 kVcm 1, 20 pulses) revealed a higher drying rate, but the rehydration rate was found subsequently to be lower (Gachovska et al., 2008; Gachovska et al., 2009). Recent investigations into the hot-air drying (70 C, 2 m s 1) of apple showed a reduction in drying time of up to 12% after the application of 50 pulses of 10 kVcm 1(Wiktor et al., 2013). However, an increase in pulse number did not necessarily lead to a higher effective moisture diffusion coefficient. Drying may also serve as a final processing step after extraction, to remove moisture from the residual plant material; an example of this is the drying of sugar beet cossettes after extraction. The PEF pretreatment of sugar beet prior to extraction has shown to improve the release of sugar, with the possibility of 208 7 Process-Induced Minimization of Mass Transfer Barriers for Improved Drying

reducing the temperature for conventional thermal cell disintegration of around 75–80 C to ambient conditions (Eshtiaghi and Knorr, 2002). The subsequently pressed pulp showed a higher dry matter content of 30% compared to 15% for the conventionally heat-extracted pulp. The higher dry matter content of the PEF pretreated sample provides significant benefits with regard to process time and energy consumption of the final hot-air drying of the cossettes. This situation is related to the water-binding capacity of the contained pectin, which is significantly increased at high disintegration and extraction temperatures. An extraction at lower temperature after PEF pretreatment results in a lower water-binding, which in turn facilitates the removal of water. This example demonstrates a secondary effect of nonthermal PEF cell disintegration in a multistep processing chain where drying occurs at a later stage and may benefit from previous processing steps being modified.

7.3.3 Impact of PEF on Freezing and Freeze-Drying Behavior of Raw Materials

Although PEF applications are targeted towards the permanent modification of cell and tissue structures based on pretreatments, beneficial effects have also been identified that result from the simultaneous application of electric fields during the freezing of noncellular material, with an impact on subsequent freeze-drying. Freezing is an operation that precedes freeze-drying, and which develops the structure (and therefore predetermines the properties) of the subsequently dried material. Freezing variables, such as the freezing rate or minimum temperature during freezing, and/or the temperature of the frozen layer during drying, can each affect the structure and drying rate of freeze-dried food and reactions in subsequent stages of a product’s life (Karel and Flink, 1973). Fast freezing rates lead to small ice crystals which form a uniform network throughout the product, leading to a dense porous matrix after sublimation. The freezing rate and ice crystal size each have direct effects on the pore size and mass flow during dehydration and reconstitution, and are therefore critical in determining the final properties of the dried product. Furthermore, in structured biological tissues, the freezing rate will also affect the location where ice crystals are formed; that is, slow cooling (below 1 K min 1)forcescrystal growth in the extracellular space, and this is promoted by the diffusion of water from the intracellular space. The existence of barriers to water movement (such as cell membranes) also needs to be taken into account. In contrast, rapid freezing at very low temperatures will produce both intracellular and extra- cellular ice crystals, with a uniform distribution and a minimum dislocation of water (Zaritzky, 2012). Especially fresh fruits, vegetables and meats undergo major changes with regard to the structural integrity of the tissue and related quality attributes after freezing, due to cell and tissue damage caused by the formation of ice crystals. Recently, advances in freezing processes have been achieved by avoiding the 7.3 Pulsed Electric Field (PEF) Application as a Pretreatment 209 formation of larger ice crystals. The acceleration of heat removal, the addition of antifreeze components and/or the development of process combinations such as high-pressure-assisted freezing or ultrasound-assisted freezing can improve the freezing process and preserve product quality (Cheftel et al., 2000; Delgado and Sun, 2011). The first step of freezing – the formation of ice crystals (nucleation) – involves the conversion of water into ice. The structure of water, with dipole moments and hydrogen bonds, as well as the attraction of water molecules and cohesive effects within liquid water, may be affected by external conditions such as the presence of an electric field, with particular impact on nucleation behavior. At the same time, extracellular ice crystal growth is related to water dislocation and osmotic moisture movement, especially during slow freezing. Pore formation in the cell membrane by the PEF pretreatment of plant tissue prior to freezing may affect this mass transfer process and the freezing characteristics, as well as water movement during subsequent sublimation. Hence, two phenomena will be discussed in the following:

The effect of an electric field applied simultaneously during freezing in order to impact nucleation and ice formation. Pulsed electric field disintegration of the cell membrane of biological tissue as a pretreatment prior to freezing, with its impact on freezing and freeze-drying characteristics.

The application of an electric field changes the orientation of the polar water molecules, and the dipole moment then aligns to the direction of the electric field vector. A higher level of structural regularity, similar to an ice-like structure, was obtained from a reorientation of the water molecules by applying an electric field (Jung et al., 1999). According to Wei et al. (2008) a critical external electrostatic field strength of 1 kVcm 1 was found to impact the orientation of the dipolar water molecules and to induce nucleation with a rise in supercooling temperature by up to 1.6 K, but no effect on phase transforma- tion time was observed. Another connection between water and electricity arises from the fact that crystal formation is associated with the occurrence of a voltage in the vicinity of the freezing front between the ice crystals and the solution (Pruppacher et al., 1968). The spontaneous growth of ice crystals in supercooled aqueous solutions is accompanied by freezing potentials of up to 15 V and freezing currents of 0.1 mA, depending on factors such as the type and concentrations of salts dissolved in the solution. Current flow in a biological tissue, and the electric disturbance that occurs during freezing by the application of an external electrostatic field, are therefore of interest for the control of crystallization phenomena. Orlowska et al. (2009) studied the effect of high-voltage electrostatic fields on ice nucleation, their aim being to control the supercooling and nucleation rates, and this in turn would result in the preservation of the microstructure of a frozen tissue. An electrostatic field with a strength of up to 60 kVcm 1 was used for this 210 7 Process-Induced Minimization of Mass Transfer Barriers for Improved Drying

purpose, and led to the induction of crystallization of supercooled water, a phenomenon called electrofreezing. The high electrostatic field was found to be capable of inducing a controlled ice nucleation with a probability of 100%. The presence of the electric field was found to modify the free energy of the system and, in consequence, the nucleation process. The nucleation rate, in turn, is related to supercooling and affects the microstructure of the frozen tissue. One different area of the application of pulsed electric fields at a lower intensity level involves pore formation in cell membranes by electroporation. This was found also to affect the freezing and freeze-drying of plant-based raw materials. When PEF was used for the pretreatment of potato by Jalte et al. (2009), an increase in PEF-induced cell damage led to an increase in the freezing rate and a reduction in the freezing times of the sample during air blast freezing. A linear correlation between the effective freezing time and the degree of cell damage was noted, but no impact on the freezing point was observed. The suggested reasons for these observations were that parts of the disrupted membranes may have served as nucleation sites that accelerated the nucleation (Toner et al., 1990), while the enhanced diffusion of extracellular and intracellular water through PEF-induced membrane pores may also have contributed to faster freezing. Furthermore, the release of cytoplasm due to membrane rupture may have influenced nucleation outside the cell and affected the freezing process due to the generated spatial concentration and temperature gradients. Freeze-drying of the samples occurred within shorter drying times for the PEF-pretreated potato such that, after a 26-h period of freeze-drying, a moisture content of only 6% was found in the sample with the highest level of cell disintegration, whereas a 16% content was detected in the control sample. Ben Ammar et al. (2010) conducted a study on the combination of PEF (cell disintegration level of 95%) and osmotic pretreatment (aqueous solution of NaCl) of potato prior to freezing and freeze-drying. The application of single pretreat- ments did not result in any visible changes of cell morphology, nor of the shape and size of starch granules in the potato tissue. However, after sequential PEF and osmotic pretreatment, a rougher surface of starch granules was observed after freeze-thawing, and this was related to the influence of electrolytes. Decreases in the freezing and freeze-drying times were observed for the PEF-pretreated as well as for the PEF- and osmotically pretreated samples compared to the untreated control samples. Noticeable structural disorders of starch granule surface morphology were demonstrated, using scanning electron microscopy, of the freeze-dried samples after sequential PEF and osmotic pretreatment. A recent study by Wu et al. (2011) investigated the effect of a PEF pretreatment on the freeze-drying rate of apple tissue. These authors concluded that the drying rate of PEF-pretreated samples was improved, and attributed this effect to a facilitated diffusion of internal water. However, no information on the repeatability of experiments and statistical significance of results was provided, which limited the validity of this study. Cryoprotectants are normally used to improve cell stability during, and cell viability after, the freeze-drying of probiotic bacteria. Complex ingredients such 7.3 Pulsed Electric Field (PEF) Application as a Pretreatment 211 as skim milk or alternative cryoprotectants such as disaccharides, polyalcohols, amino acids and proteins have been found to be suitable (Siaterlis et al., 2009). When freezing and freeze-drying of bacteria, suspensions are used in which the cryoprotectant can be dissolved and interact with the suspended single microbial cell. For complex vegetable tissues, the application of cryoprotectants is at first limited to intrinsic compounds as the transfer of external cryoprotectants into the cells is hindered. However, cryoprotection can serve as an efficient tool to improve the quality of frozen products. The accumulation of osmotically active substances in the cytosol of the cell forms part of this protective mechanism. Based on their protective and stabilizing effects in terms of decreasing the chemical potential of water and the freezing point in the cytoplasm, as well as improving the cell membrane stability, sugars may be used as cryoprotectants. Indeed, industrial techniques have been developed (such as vacuum impregnation) in order to fill the pores of plant tissue with an external cryoprotectant, the transfer of which into the cell could further improve its efficiency, though this is limited by the cell membrane acting as a mass transfer barrier. An innovative concept to improve the freezing tolerance of plant tissue by the vacuum impregnation of trehalose in combination with a PEF treatment was investigated by Phoon et al. (2008). In this case, PEF was applied to modify the plasma membrane organization, followedbyalocalandtemporarypermeabili- zation in order to avoid permanent tissue damage. Such a process was used to transport trehalose across the cell membrane, leading to an infusion into the cells; the result was that, in combination with vacuum impregnation, the freezing resistance of spinach leaves was improved. Further investigations are required to evaluate the effect of PEF pretreatment on the performance of cryoprotectants in the freeze-drying of complex tissues and the resulting product quality characteristics.

7.3.4 Quality Characteristics Affected by PEF Pretreatment

Parameters such as cell disintegration index, porosity, water content and distribu- tion, water activity, color, texture, shrinkage, rehydration behavior and leakage, content of heat-sensitive compounds and distribution of dissolved solids, as well as overall sensorial perception, are of interest when evaluating the impact of PEF and ultrasound, and also of the drying process, on product quality characteristics. The physical and biochemical status of foods is affected by conventional dehydration methods, which contribute to changes in the color, texture and taste of heat-treated products, as well as to the loss of valuable heat-sensitive components. Differences in shrinkage during drying depend on the raw material, and have been found to be more pronounced for higher-porosity products such as apple compared to potato or carrot (Ratti, 1994). Differences in texture and in drying, as well as rehydration behavior, are also based on the initial water content (with average values of 86% for apple, 91% for carrot and 80% for potato), as well as on 212 7 Process-Induced Minimization of Mass Transfer Barriers for Improved Drying

the water-holding capacity (12 g g 1 for apple, 23 g g 1 for carrot, 2 g g 1 for potato) (McConnell et al., 1974). Chemical changes such as degradation or oxidation reactions are undesired, and depend not only on the local temperatures that are reached during drying but also on the local water activity levels achieved during and after drying (Miskin et al., 1983). It is not the average water content that determines such reactions, but the distribution of free water after drying with local maximum values that contribute to chemical or microbial quality degradation (Wang and Brennan, 1995). Hence, a faster and more homogeneous drying resulting in a more uniform distribution of residual moisture due to enhanced mass transfer, especially in larger pieces, might be more beneficial with regards to product quality and shelf stability. The use of PEF as a pretreatment showed a significant effect on some of the quality parameters studied. Taiwo et al. (2001) investigated the impact of different pretreatments for cell disintegration such as blanching, freezing, high pressure and PEF on the vitamin C content of apples after osmotic dehydration. During the first 2 h of dehydration, the PEF-pretreated samples retained the highest vitamin C levels compared to the other pretreatments. However, during further osmotic dehydration of up to 6 h, the highest vitamin C retention of around 10% of the initial value occurred in the untreated and high-pressure-treated samples. The greater loss of vitamin C in PEF-pretreated samples was attributed to a faster leaching as a result of the cells being permeabilized. However, a minimal depletion of b-carotene has been reported for osmotically dehydrated PEF-pretreated red bell pepper samples (Ade-Omowaye et al., 2002). Taiwo et al. (2002) reported a substantial increase in rehydration ability with regards to rehydration time and rehydration capacity after the osmotic dehydration of PEF-pretreated apple slices. Osmotic dehydration itself showed a higher rate, with a reduction in process duration by up to 40%. The kinetics of water gain and solids loss during the rehydration of untreated and PEF-pretreated carrot samples after osmotic dehydration were studied by Amami et al. (2007). PEF pretreatment was found to increase the rehydration capacity of osmotically dehydrated samples and reduce the solids loss. A higher retention of solids during the rehydration of osmotically dehydrated samples was also reported by Taiwo et al. (2002) for apples. PEF pretreatment improved the rehydration capacity of carrots in combination with osmotic dehydration; however, PEF samples subjected to drying (hot-air drying at 80 C) without previous osmotic dehydration showed the lowest rehydration capacity. This effect was attributed to the higher solid uptake achieved for PEF-pretreated samples during osmotic dehydration, which results in a minimized shrinkage (Amami et al., 2007). In the same study, PEF pretreatment was found to slightly decrease the firmness of rehydrated samples, as determined by measurements of the relaxation curves of rehydrated carrot samples. Bazhal et al. (2003) investigated the shrinkage and porosity values of apple after hot-air drying, and found a more severe collapse of structure in the control samples compared to PEF-treated samples. The bulk and skeletal density values were 0.46 7.3 Pulsed Electric Field (PEF) Application as a Pretreatment 213

and 1.39 kg m 3 for the control samples, respectively, and 0.38 and 1.52 kg m 3 for the PEF-pretreated samples. The PEF treatment of samples resulted in lower bulk densities due to the resultant less shrinkage during drying of the treated materials. The higher mean skeletal density for the PEF-pretreated samples was attributed to the increased porosity resulting from electroporation. With regards to the impact of PEF on freezing and freeze-drying (see Section 7.3.3), it was noted that an enhanced diffusion of extracellular and intracellular water would lead to smaller sizes of formed ice crystals, and to a better quality of product. However, even if small ice crystals were formed during the freezing process, they may grow larger during storage of the frozen food due to recrystallization; this could result in additional membrane damage and reduced water-holding capacity if not followed by freeze-drying within a certain time. In the study of Jalte et al. (2009) on the freezing and freeze-drying of PEF- pretreated potato, the samples were reported to be more uniform in shape with clearer colors, a smaller degree of shrinkage and a lower browning level. The rehydration characteristics, determined by the leakage of ionic components in the solution, were similar between the PEF-pretreated and untreated samples for the first 50 min in distilled water at 25 C, but the release of ionic compounds was subsequently increased for the PEF-pretreated sample. The use of a combination of PEF and osmotic pretreatment prior to freezing and freeze-drying caused not only a decrease in the freezing and freeze-drying times but also a visually better quality of the pretreated and dried samples, as reported by Ben Ammar et al. (2010). As PEF affects water transport, the distribution of soluble solids might also be affected during and after the drying process, with an impact on product quality and sensorial attributes. An investigation into the water transport, distribution of soluble solids, and the location of the drying front was conducted on the tissue level by J€ager et al. (2010) during the convective hot-air drying of apple cubes. An analysis of the drying curves showed a 21% reduction in drying time in case of the PEF-pretreated samples in comparison to control samples in order to reach the common moisture level of industrially dried apple products of 20% (wet basis). Higher drying rates were observed for the PEF-treated samples, and the drying front was found to migrate faster towards the core of the sample, as analyzed using magnetic resonance imaging (MRI). A lower standard deviation for the moisture content at different drying times was obtained for the PEF-treated samples, indicating a more homogeneous drying process as a result of the PEF-induced cell disintegration which compensates textural differences in the raw material properties. The migration of fructose, depending on drying time and pretreatment, is shown in Fig. 7.7, which illustrates the fructose ratio between the core and the surface of the apple cube (values above unity indicate a higher fructose concentration in the core). Fructose distribution in the sample was found to differ significantly between the PEF-pretreated and untreated apple cubes. In untreated samples, the fructose ratio was decreased, and showed an increasing concentration of fructose in the surface layer as drying proceeded. This finding was expected, as the dissolved 214 7 Process-Induced Minimization of Mass Transfer Barriers for Improved Drying

Fig. 7.7 Effect of PEF pretreatment on calculated based on dry matter at the fructose distribution in apple cubes (core corresponding drying time. Reproduced with and surface layer) during hot-air drying. The permission from J€ager et al. (2010). fructose concentration of the two parts was

fructose migrates and changes its location during drying. Moisture (liquid rather than water vapor) moves towards the surface of the apple cube such that, as the solution nears the surface pure water evaporates from it, leaving behind an increasingly concentrated solution in the outer layer. In the completely dry product, a gradient of soluble material should be apparent with the highest concentration at the surface of the apple cube. However, fructose migration in the PEF-treated sample showed an opposite behavior, with a fructose ratio of more than unity, indicating higher fructose concentrations in the core of the apple cube. This migration is also more pronounced when the drying time is extended. The water distribution analyzed by MRI in the control and PEF-treated apple cube after 2 h of drying is illustrated in Fig 7.8. The PEF-pretreated samples showed a higher level of shrinkage at the same drying time, since higher water losses had occurred in comparison to the control sample. In order to characterize the location of the drying front, the relative water content, given as the averaged nuclear magnetic resonance signal intensity within a rectangular region of interest, is shown. The control sample shows a lower water content within a 3 mm surface layer, and a larger high moisture zone within the core of the apple cube. For the PEF-treated sample, the surface layer with a lower water content has a dimension in the range of 4–6 mm, indicating that the drying front is already located closer to the center of the cube. When the wetted area on the surface has totally disappeared, the liquid recedes into the capillaries and moves deeper as drying continues. This occurs more rapidly for the PEF-treated sample than for the control sample. The 7.3 Pulsed Electric Field (PEF) Application as a Pretreatment 215

Fig. 7.8 (a) Water distribution in the middle Water distribution along the center line of layer of the apple cube (top, control; the apple cube, as determined using MRI. bottom, PEF) after 2 h of drying. The white Reproduced with permission from J€ager color indicates a higher water content; (b) et al. (2010).

evaporation occurs at a certain distance below the food surface, and the diffusion of vapor occurs from the site of vaporization to the surface. Since, for the PEF-pretreated sample the drying front moves faster and is thus located closer to the center of the apple cube, the evaporation of water occurs in the deeper layers within the apple cube. Water vapor diffuses through pores, cracks and capillaries of the tissue, and the evaporation of water in the deeper layers leaves behind an increasingly concentrated solution in the core of the apple cube, which may result in a higher fructose concentration. Hence, the lower fructose concentration found in the surface layer of the cube could be partly related to differences in the water transport in PEF-permeabilized samples. This example shows that cell disintegration achieved by PEF does not only affect the drying process in terms of facilitating water removal. In addition, the modification of water transport also affects the distribution of solutes such as fructose. Despite the reduction in drying times, these results exemplify the complex interactions between the PEF process and cell disintegration, the modification of structure, and the related mass transfer processes and product functional attributes such as sugar content distribution. Furthermore, the data obtained on fructose distribution reveals the possibility of a tailor-made modification of final product characteristics. The ability to affect the distribution of soluble compounds may provide the potential to increase the storage stabilities of light- or oxygen-sensitive constituents by transferring them to the center of the product, and also to modify various sensorial product properties. Investigations into the enhancement of sweetness intensity in gels by the 216 7 Process-Induced Minimization of Mass Transfer Barriers for Improved Drying

inhomogeneous distribution of sucrose have revealed that samples with an inhomogeneous sucrose concentration were perceived as being sweeter than samples with the same content of sucrose but with a homogeneous distribution. As a consequence, an inhomogeneous distribution of sucrose can be used to reduce the sucrose content of a product by 20%, without any decrease in sweetness intensity (Mosca et al., 2010, 2012a, 2012b).

7.4 Contact Ultrasound for Combined Drying Processes

Ultrasound-assisted drying processes can be based on sound transmission in air (airborne ultrasound), liquids (osmotic solutions) or solids (contact ultrasound). Sound transmission and sound emitter designs must be adapted to the respective applications and have, together with the transmission medium, a major influence on process design and performance. Air is the natural transmission medium in most drying processes. Challenges occur due to the poor impedance matching between ultrasound emitters and air, and the easy disruption of the sound field by high air velocities (Gallego-Juarez, 1998). Successful process design has been reported using specially designed stepped-plate transducers that allow a minimization of impedance mismatch (Gallego-Juarez et al., 1994), as well as using vibrating drying chambers in fluidized-bed drying (García-Perez et al., 2006). However, limitations persist in terms of air velocity and, most likely, also scale-up. In addition, air-drying or fluidized-bed drying are not suitable for all products due to problems of heat sensitivity or special requirements for product characteristics; for example, a high rehydration capacity is only achievable by freeze-drying. Sound transfer in liquids has been well studied, and ultrasonic baths or probe systems have been used to transfer ultrasonic waves to water or osmotic solutions in order to increase mass transfer and disintegrate cells by cavitation effects. The use of liquids as transmitting media is suitable for dehydration processes when osmotic solutions, such as sugar or salt solutions, are used. The mass transfer- enhancing effect of ultrasonic waves increases the exchange of the treatment medium and product, and can thus increase the osmotic drying effect (Simal et al., 1998; Deng and Zhao, 2008). However, there may be a consequential leakage and uptake of sugars or salts, while the loss of aroma, flavor and color, as well as a need to declare sugar or salt contents on the food packaging, represent reasons to seek alternatives. Contact ultrasound is of special interest for use in drying processes as it allows the product to be affected without any need for an additional sound-transmission medium. The high water content of food products, especially at the start of a dehydration process, allows good impedance matching and improved sound transmission. Contact ultrasound-assisted drying employs transducers placed in contact with the product (Gallego-Juarez, 1998), or it may use the support of the products directly as the sound-transmitting surface (Schossler€ et al., 2012a, 2012b). 7.4 Contact Ultrasound for Combined Drying Processes 217

Ultrasonic waves can be excited into screens, which transfer the vibration to the food product, and the ultrasound emitter can be attached to the center of the screen, as well as to its frame. Direct transmission to the product allows sonication to be used in a large variety of food processes, throughout the complete drying process. Future equipment developments could enable ultrasound-mediated transfer not only to screens but also to trays or conveyor belts, and in this way air- drying in batch systems, as well as in continuous systems, could be assisted by ultrasonic vibrations. Even in vacuum freeze-drying, where air transmission is impossible due to the vacuum, the product can be placed on supports and treated with ultrasonic waves (Schossler€ et al., 2012b).

7.4.1 Ultrasound in Osmotic Dehydration

Although the application of ultrasound in osmotic dehydration is, technically speaking, not a contact ultrasound application, its analysis provides an interesting insight into the principles of ultrasound-assisted drying and the different approaches to minimize mass transfer barriers applying PEF and ultrasound. It is therefore briefly discussed here. In terms of equipment design, ultrasound-assisted osmotic dehydration can benefit from developments made in the ultrasonic processing of liquids. Ultrasonic baths or probe systems can be used, while multiple ultrasound processors can be connected in series in order to enable the uniform sonication of large volumes. Sound transmission from sonotrode to liquid is well mastered by existing equipment. Cavitation will occur when power ultrasound is transmitted in liquids, and this adds an important benefit to the range of ultrasound effects improving mass transfer. In addition to the cell-disrupting action of cavitation, which is important in liquid applications, mechanical influences such as microstreaming, shear, liquid jets and associated shock waves can directly impact mass transfer and boundary layers. Furthermore, the so-called hydrodynamic mechanism of mass transfer, which describes the improved penetration with osmotic solution after the removal of gas from the pores of plant tissue (Fito and Pastor, 1994), has been assumed to be part of the ultrasound-related mechanisms improving osmotic dehydration (Simal et al., 1998). In the case of ultrasound-assisted osmosis, investigations have been primarily focused on the removal of water from fruits by using hypertonic sugar solutions, mainly based on sucrose. Cárcel et al. (2002) investigated the osmotic dehydration of apple in 30 Bx sucrose solution by comparing ultrasound at intensities from 12.5 to 16.6 Wcm 2 to static conditions, as well as agitation at 850 and 1200 rpm. These authors showed that ultrasound only affected water loss and sugar gain above a certain threshold of ultrasonic intensity, found in this study to be in the range of 14 to 16 Wcm 2. This result was comparable to the effects of ultrasound observed in meat brining, where the intensity applied during ultrasonic treatment was required to be above a threshold of 39 Wcm 2 to influence the moisture content, and above 51 Wcm 2 to influence the NaCl content (Carcel et al., 2007). 218 7 Process-Induced Minimization of Mass Transfer Barriers for Improved Drying

In addition to direct ultrasound application, successful ultrasound pretreatments leading to an increase in drying rates for subsequent air-drying were reported for fruits and vegetables. In these studies distilled water was used as a treatment medium, in addition to hypertonic sugar or salt solutions. Distilled water itself is an osmotically active liquid, which will not cause the removal of water from the cell; rather, water uptake will occur, reducing the osmotic difference between cell interior and surrounding medium. This leads to a bloating of cells, as observed in a study on the sonication of melon cubes (Fernandes et al., 2008a). Furthermore, electrolyte loss may occur, as reported for the ultrasonic pretreatment of mush- rooms, cauliflower and Brussels sprouts (Rezek Jambrak et al., 2007). These effects can impair the structure of the cell membrane and weaken its resistance against water removal from the cell. This assumption is in line with observations made in other studies (Fernandes et al., 2008c; Fernandes et al., 2008b), in which improved drying was reported not only for pineapple and papaya samples subjected to sonication but also for samples that had just been allowed to soak in distilled water. The application of ultrasound seemed to improve penetration with distilled water, increasing the effect on the plant cells. It must be assumed, that the effects of ultrasound depend on the structure and porosity of the treated product (as discussed above), and that the penetration of dense structures as potato tissue would probably be impaired. However, it would appear that no investigations on this topic have yet been reported.

7.4.2 Contact Ultrasound in Air Drying

In order to overcome the limitations represented by the disturbance of the sound field at high air velocities, Gallego-Juarez (1998) tested an improvement in the application of a stepped-plate transducer, whereby a vibrating plate was lowered until it came into contact with the samples. The experiments revealed improved drying characteristics compared to conventional hot-air drying and to the application of airborne ultrasound, and even permitted a reduction in the final moisture content of carrot samples to less than 1%. The most important ultrasonic effects leading to these improved drying characteristics were considered to be the sponge effect and the creation of dewatering channels by acoustic stress. In an associated study, an improved dehydration of carrots and mushrooms was achieved either by placing the products on top of the transducer or by placing them below the transducer and applying a static pressure to the support plate (Riera-Franco de Sarabia et al., 2002). An associated patent using ultrasonic diffusing plates in contact with the product as it is transported on a conveyor belt (Gallego-Juarez et al., 2001) brought the technology of contact ultrasound in hot-air drying to an industrial level. In this case, the product is transported on a conveyor belt and dried using a hot gas stream. A static pressure exerted from the top side of the ultrasonic diffusing plates assures that contact between the product and the ultrasound device is maintained throughout the drying process. 7.4 Contact Ultrasound for Combined Drying Processes 219

Contact ultrasound-assisted drying using screens as both sample-supporting and sound-transmitting surfaces was described in studies devised to analyze the enhancement of bell pepper, apple and potato drying (Schossler€ et al., 2012a, 2012c). These studies presented a solution for adapting contact ultrasound-assisted drying to a large range of industrial drying processes, with the systems being based on the industrially well-established technology of ultrasound-assisted screening (Hidaka and Miwa, 1979; Monteith, 1992; Peters, 1996; Obradovic and Brnci9 c, 2009; Prokic, 2010). Although the ultrasound transducers were adapted to the vibration characteristics of the screen frames, the screens themselves did not have to be specifically adapted in order to allow successful sound transmission to the product. The sonication of larger screens can, for example, be realized by using waveguides in order to increase the area of sound transmission. One interesting advantage of screens compared to closed trays is their larger open surface for water removal. The intensity of the ultrasound treatments is characterized by the excitation amplitude at the contact point of sonotrode and screen frame, and can be measured using laser interferometry. This excitation amplitude is increased on the screen mesh by factor of three to six, depending on the system used (Schossler€ et al., 2012a, 2012b). The amplification of the ultrasonic wave at the treatment surface allows a reduction in excitation amplitude at the frame, and thus a reduction in energy consumption of the ultrasound generator. Sound transfer to a standard laboratory screen with an aperture of 500 mm used in the studies allowed the air-drying time of apples to be reduced by 27%, and of bell peppers by 23% (Schossler€ et al., 2012a). The ultrasound system used is shown in Fig. 7.9, and the corresponding drying curves are depicted in Fig. 7.10; these show the reduction of the moisture ratio MR over the drying

Fig. 7.9 Assembly of the ultrasonic system used in contact ultrasound-assisted air drying. 1, samples; 2, drying screen; 3, ring sonotrode RIS 200; 4, ultrasound processor UIS250L. Reproduced with permission from Schossler€ et al. (2012a). 220 7 Process-Induced Minimization of Mass Transfer Barriers for Improved Drying

Fig. 7.10 Drying curves of apple and bell analysis of significance (t-test) with different pepper cubes dried conventionally (Control) letters at the same drying time representing and with the application of continuous statistically significant differences of the ultrasound (US) with lines representing the moisture ratio (MR) value at a confidence calculated drying curves for the respective level of 95%. Only values at the same drying control treatments according to Midilli et al. time are compared. Reproduced with (2002). The table presents the results of the permission from Schossler€ et al. (2012a).

time. The dimensionless moisture ratio is defined as X X MR ¼ eq ; ð7:4Þ X0 Xeq

where X is the moisture content at the given time, Xeq is the equilibrium moisture content, and X0 is the initial moisture content (all in kgwater per kgdry solids). When the relative air humidity during a drying process is not controlled and the equilibrium moisture content cannot be determined, MR can be simplified to ¼ € MR X/X0 (Schossler et al., 2012a). The MR of both apple and bell pepper was significantly reduced by the application of continuous ultrasound from the first measurement until the final moisture content was reached. At the same time, sample shrinkage was not affected by the ultrasound waves, indicating a gentle improvement of the water removal. Further reductions in the energy consumption of contact ultrasound-assisted drying can be realized by applying an intermittent ultrasound treatment. Whereas, the maximum impact of ultrasound is achieved when the sound waves are 7.4 Contact Ultrasound for Combined Drying Processes 221 continuously transmitted into the product, optimization between the drying time reduction and additional energy consumption by the ultrasound system is essential for a successful industrial application of the technology. The net sonication time applied during a drying process can be reduced applying interval switching of the ultrasound system. In experiments conducted on apple drying, the application of continuous ultrasound at an excitation amplitude of 4 mm led to a reduction in drying time by 27%, and a reduction by 20% was achieved applying ultrasound in an interval of 10 s sonication and 10 s recovery time (Schossler€ et al., 2012a). Hence, a net sonication time of 50% allowed about 75% of the ultrasound effect to be maintained. There does, however, seem to be a certain threshold below which an intermittent ultrasound application can no longer significantly affect the drying. In the above-cited study, no impact of intermittent ultrasound was found when the net sonication time was reduced to 10% by applying 10 s of sonication followed by 90 s recovery time. In order to allow economic drying processes assisted by contact ultrasound, the relationships between sonication time, intensity and drying enhancement must be optimized for the product of interest. However, the large potential of this technology has already been demonstrated by not only results obtained at the laboratory scale but also the large product range tested. Significant effects on drying times were achieved for tissues ranging from highly porous (apple) (Schossler€ et al., 2012a) to practically nonporous tissue (potato) (Schossler€ et al., 2012c). Bell peppers represent a product with special drying characteristics, as the peel is not removed prior to drying (Schossler€ et al., 2012a). The exicarp of the bell pepper consists of epidermal cells with thickened cell walls (Gassner et al., 1989) that protect the plant tissues against environmental stresses and moisture losses, with moisture transfer being much-reduced from the peel-covered side of the tissue. Hence, in order to closely represent industrial conditions, where the product bulk would be placed randomly in a single layer on the drying screen or tray, bell pepper samples subjected to hot-air drying were not arranged on the screen according to the location of the peel (Schossler€ et al., 2012a). The application of contact ultrasound-enhanced air drying of the above- mentioned products has led to a positive evaluation of the industrial applicability of this technology. The application of a static pressure in order to further improve sound transmission, as described earlier (Gallego-Juarez, 1998; Riera-Franco de Sarabia et al., 2002), represents an interesting route for future process optimiza- tion, although equipment still needs to be improved and updated at the industrial scale.

7.4.3 Contact Ultrasound in Freeze-Drying

Freeze-drying is a food-conservation process that leads to unique product characteristics. The removal of water from a frozen product under vacuum conditions results in a near-perfect preservation of shape and color, and leads to a sponge-like structure that allows a very fast water uptake during rehydration and 222 7 Process-Induced Minimization of Mass Transfer Barriers for Improved Drying

characteristics of freshness to be regained (Barbosa-Canovas and Vega-Mercado, 1996). The importance of the freezing step itself has been discussed in detail in Section 7.3.3. The creation of small ice crystals is of special interest in terms of good product quality, and not only PEF pretreatments but also ultrasound-assisted freezing can be applied to assure the formation of small ice crystals with a narrow size distribution (Acton and Morris, 1992). In general, the two most important effects are an ultrasound-induced enhancement of heat and mass transfer, and the fracture of ice crystals by cavitation, leading to smaller ice crystals. Zheng and Sun (2006) presented a detailed review on the application of ultrasound for freezing foods that is not only limited to the optimization of crystal size distribution; rather, the freezing time may also be shortened and there may be benefits for partial freezing with surface freezers, so as to minimize fouling. In vacuum freeze-drying the removal of water due to sublimation is relatively slow, such that the production capacity of freeze-drying plants is limited by long drying times of several hours. In general, the process consists of two drying stages: during primary drying, ice is sublimated under reduced pressure conditions, whereas secondary drying is aimed at evaporating the bound water. In industrial freeze-drying, both drying stages are assisted by an external heating source (mostly heated trays or infrared radiation), which provides the energy for the sublimation and evaporation. However, these systems are limited with regards to the temperature of the heating medium in order to avoid product damage due to overheating. The high energy consumption and long duration of vacuum freeze-drying makes it a very interesting process for technologies that allow drying time reductions; however, the transfer of sound waves under vacuum is greatly limited. Early studies on the nonvacuum freeze-drying of tea and coffee extracts using airborne ultrasound and a nozzle-based system showed the effectiveness of ultrasonic waves in the enhanced drying of frozen products (Moy and DiMarco, 1970). An increase in the mass transfer coefficient, a reduction of vapor pressure in the air-stream during the rarefaction cycle, a thermal effect based on friction and adiabatic compression, and the sound pressure energy were each discussed as possible mechanisms behind the observed increases in freeze-drying rates. In a more recent study, this principle was transferred to the drying of tissue-based foods when Bantle et al. (2010) dried ice and several food products in an ultrasound-assisted atmospheric freeze-drying process and observed increases in the sublimation rate of 23.3% for ice and 6% to 17.5% for food products. The system devised by these authors consisted of a modified fluidized-bed drying chamber, where an ultrasound transducer originally designed for welding purposes was placed beneath the bottom plate. The freeze-drying experiments were then carried out with stationary as well as fluidized beds. Although promising results were achieved with atmospheric freeze-drying, the application of airborne ultrasound is in this case limited in the same way as in hot- air drying. The low impedance matching between air and the ultrasound transducer must be overcome by special transducer design, but high air-velocities 7.4 Contact Ultrasound for Combined Drying Processes 223 in fluidized beds may disrupt the ultrasonic field. In the food industry, a wide variety of sensitive products can be dried using tray-based vacuum freeze-drying processes. Contact ultrasound represents an interesting means of employing the drying-enhancing effects of ultrasound under low-pressure conditions. As with contact ultrasound-assisted air drying, the ultrasound transducer can be coupled directly to the product support or its frame. This design would be easily adaptable to today’s industrial reality in the freeze-drying of foods (Schossler€ et al., 2012b). Unfortunately, the low-pressure conditions in vacuum freeze-drying may lead to certain challenges in terms of ultrasound application. Notably, part of the ultrasonic energy transmitted to the product is transformed into heat due to interparticle friction, scattering and viscous damping (Ensminger, 1988), and the sample temperature is crucial to the quality of the freeze-dried products. An increase in product temperature above the glass transition temperature can lead to ice melting and consequential structural collapse and chemical alterations (Franks, 1998). As drying progresses, however, the risk of overheating is increased as heat transfer to the surrounding air is impaired under low-pressure conditions. In terms of system design, certain challenges exist regarding the integration of an ultrasound transducer into the vacuum chamber. The transformation of electrical into mechanical energy (namely the ultrasonic vibrations) leads to heat generation in the transducer. Under ambient conditions, transducers can dissipate the generated heat energy to the surrounding air, thus avoiding any overheating of the equipment. At the laboratory scale, the impaired heat dissipation in the vacuum chamber can be avoided by installing the transducer outside the vacuum chamber. A system for contact ultrasound-assisted vacuum freeze-drying at the laboratory scale is shown in Fig. 7.11 (Schossler€ et al., 2012b) where, in order to transmit the ultrasonic vibrations to a sample-supporting screen in the vacuum chamber, a sonotrode was coupled to the transducer and passed through the lid of the vacuum chamber with a vibration-free flange. The study by Schossler€ et al. (2012b) analyzed the impact of ultrasound parameters on sample heating and drying improvements of bell peppers in relation to product quality. Experiments on the impact of ultrasound amplitude and the applied interval during intermittent sonication showed that sample overheating and thawing can easily occur, and that these parameters must be controlled accurately. In contrast, the results also showed that a correct choice of parameters (amplitude, interval of ultrasound application) would enable contact ultrasound- assisted freeze-drying with controlled sample heating and maintained product quality, as confirmed by analyses in terms of bulk density (volume retention), color, ascorbic acid content and the rehydration characteristics of dried bell pepper cubes. An interval application with 10% net sonication time, realized as 10 s sonication and 90 s recovery time, led to a drying time reduction of 11.5%. Figure 7.12 shows the corresponding drying and temperature curves. Bell pepper cubes dried with the application of contact ultrasound showed a significantly faster reduction of the MR than did those dried without ultrasound assistance. The temperature curves showed no additional heating of the ultrasound-treated samples for 7 h of drying, which indicated that the sound energy had been used directly for sublimation. On 224 7 Process-Induced Minimization of Mass Transfer Barriers for Improved Drying

Fig. 7.11 Laboratory-scale system for 4, acrylic lid; 5, ultrasound drying screen; 6, contact ultrasound-assisted freeze-drying. A, thermocouples; 7, drying chamber; 8, rack; Acrylic lid with ultrasonic system; B, screen 9, freeze-dryer; 10, water outlet; 11, vacuum for control samples; C, freeze-dryer; 1, regulation; 12, vacuum tube; 13, display. ultrasound processor UIP1000; 2, Reproduced with permission from Schossler€ sonotrodes BS2d34; 3, vibration-free flange; et al. (2012c).

the completion of primary drying (after 7 h) the temperature of the ultrasound- treated samples was increased, whilst during secondary drying the sound energy was used to evaporate bound water such that thermal effects occurred. The control samples reached the surrounding temperature of 24.2 C after 18.2 h of drying, whereas samples dried with the application of contact ultrasound reached this value after only 16.2 h of drying. This reduction in drying time of 2 h was confirmed by mathematical modeling of the MR curves (Schossler€ et al., 2012b). The higher final temperature of the ultrasound-treated samples was attributed to the conversion of sound energy to thermal energy when the final moisture content was reached, and the sound energy was no longer used for evaporation.

7.4.4 Quality Characteristics Affected by Ultrasound-Combined Drying Processes

Ultrasound influences the product via a variety of mechanical, thermal, and chemical effects. The mechanical effects are mainly related to the compression and 7.4 Contact Ultrasound for Combined Drying Processes 225

Fig. 7.12 Effect of intermittent ultrasound bell pepper in comparison to conventionally treatment with 10 s sonication and 90 s dried samples (non-US). Reproduced with recovery (US) on moisture ratio (MR) and permission from Schossler€ et al. (2012b). sample temperature during freeze-drying of rarefaction caused by the sound wave. When acting on a liquid medium, these effects can lead to cavitation with associated phenomena, such as shear forces, microstreaming, liquid jets and shockwaves. They can also create a strong mechanical impact on a product present in the sound field (Suslick, 2003). Due to the extreme conditions that exist during bubble collapse, with local temperatures and pressure peaks, chemical changes can occur and even lead to radical formation. These effects are especially important for ultrasound treatment in liquids, and deterioration has been reported due to the oxidation of fats, oils, vitamins, or color pigments (Roberts, 1993). As a consequence, quality impairment has been reported for certain products subjected to ultrasound-assisted osmotic dehydration. An increased cell disruption and a greater loss of firmness were observed in apple samples (Deng and Zhao, 2008), while Stojanovic and Silva (2007) demonstrated an increased loss of phytonutrients in rabbiteye blueberries after the application of ultrasonically assisted osmoconcentration prior to hot-air drying, when compared to conventional osmoconcentration. In contrast, contact as well as airborne ultrasound application avoids the occurrence of external cavitation and thus reduces the risk of related effects such as radical formation and oxidation. Thus, to the present author’s knowledge, no important quality impairment due to such ultrasound-assisted processes has yet been reported. Soria et al. (2010) found only minor changes in the reducing sugar 226 7 Process-Induced Minimization of Mass Transfer Barriers for Improved Drying

content, antioxidant activity and protein profiles of blanched carrots subjected to airborne ultrasonic drying; instead, the rehydration characteristics were even improved. Studies conducted on heating effects in contact ultrasound-assisted vacuum freeze-drying showed that the parameters of such ultrasound applications must be well adapted in order to avoid product damage (Schossler€ et al., 2012b). However, the results of this study also proved that it is possible to design processes with important improvements in drying rates without causing a negative effect on product quality in terms of shrinkage, color, rehydration characteristics or the content of compounds, such as ascorbic acid. These findings were in line with the results of a study on ultrasound-assisted atmospheric freeze-drying of green peas, in which ultrasound treatment had no impact on color and shrinkage of the product (Bantle and Eikevik, 2011). The fact that the mechanism of action of ultrasound in tissues seems to be primarily mechanical, and perhaps partly thermal at the tissue level (Schossler€ et al., 2011; Schossler€ et al., 2012c), may explain this limited effect on product quality. Yet, as long as the harsh conditions of cavitation can be avoided the product quality can be maintained when the parameters of the ultrasound treatment are well chosen. Nonetheless, detailed investigations of product quality must accom- pany the optimization of process parameters, especially in the case of critical applications, such as ultrasound-assisted osmotic dehydration, or applications involving special conditions, such as vacuum freeze-drying.

7.5 Conclusion

The underlying mechanisms of the modifications of cell and tissue structures by PEF and ultrasound, and their related effects on drying, have not yet been fully elucidated and, in most cases, are still limited to specific raw materials and drying processes. An optimization of existing systems, and the design of new processing concepts, require further investigation in order to establish general models that will reduce experimental trials and allow a transfer of results to other raw materials and processing parameters. For this purpose, advanced analytical methods such as MRI will be required in order to reveal particularities in heat and mass transfer, as well as structural modifications. Nonetheless, numerous studies have proved that the cell disintegration caused by PEFs, when applied to plant tissues, has a related effect on mass transfer during osmotic dehydration, air-drying or freeze-drying, with drying rates possibly being increased for all such drying processes by PEF pretreatment of the raw material. With most previous investigations having focused on osmotic dehydration and convective air drying, there are to date only a small number of studies available on the modification of freeze-drying characteristics of plant tissues subjected to PEF pretreatment before freezing. Further investigations are required in order to establish any clear correlation between the degree of cell disintegration by PEF and its impact on ice crystal formation, as well as on water removal during sublimation. Different tissue structures 7.5 Conclusion 227 must be taken into account in order to establish a link between drying characteristics and structural particularities. Concepts related to PEF-mediated cell disintegration that have been used to improve the infusion of cryoprotec- tants require further investigation also with regard to the impact not only on freezing but also on freeze-drying behavior and final product quality. The impact on nucleation of electrostatic fields applied during freezing was clearly demonstrated for pure water and for water containing solutes. Hence, investigations need to be extended towards more complex liquids as well as to tissue structures in order to study any possible modifications of supercooling and nucleation rates, and to determine the resultant impact on the microstructure and quality of tissues. No concise conclusion can yet be drawn regarding the effect of PEF pretreatment on product quality after drying, as the amount of data currently available is very limited. Shrinkage was reduced, and rehydration ability improved, for PEF-pretreated samples in most cases, in addition to a visually better quality. However, losses of vitamin C during osmotic dehydration were more pronounced for PEF-pretreated samples than for untreated samples. Modifications of the cell structure by PEF showed effects not only on water removal but also on the dislocation of dissolved solids, as noted by the changes in sugar distribution in the case of hot-air drying of apple cubes. The latest studies analyzing the mechanisms of contact ultrasound-assisted mass transfer have revealed that ultrasound-related cell disruption is limited to the thin layer of tissue that is in direct contact with the vibrating surface. In any tissue layers deeper than 1 mm, the structural changes that occur are attributed to water removal; moreover, ultrasound can enhance such water removal as a result of mechanical effects that can loosen cell–cell bonding, cause microstreaming, reduce the thickness of boundary layers, and remove moisture from solid–liquid interfaces. The transport of moisture through the pore network can also be increased, and low- pressure cycles will enable an improved evaporation in the tissues. The mechanical effects that act at tissue level generally lead to a good retention of product quality, and until now many of the studies describing the quality of food products dried by the application of contact ultrasound have failed to reveal any negative effects. However, many of the studies cited in this chapter have shown the importance of a good adaptation of process parameters, especially when processes such as vacuum freeze-drying are involved. The thermal effects associated with ultrasound transmission are especially impor- tant in a low-pressure environment, where heat exchange with the surroundings is impaired. Strong sample heating during the primary drying can also lead to melting and structural collapse. Parameter optimization in efficient processes is much more difficult when ultrasound is transmitted using liquids, as in the case of ultrasound-assisted osmotic dehydration. The strong physical and chemical effects of cavitation can easily cause cell disruption, with an associated loss of tissue firmness, leakage, and oxidation. The implementation of PEF treatments in existing industrial processing lines has been realized for numerous installations, based not only on adapted treatment chamber designs but also on the availability of industrial-scale pulse modulators. Although the integration of PEF treatments into complex processing 228 7 Process-Induced Minimization of Mass Transfer Barriers for Improved Drying

concepts is feasible from a technical point of view, it is still important to consider the performance of any connected processing steps. Indeed, the consideration of textural properties and the particle and cell sizes of the different raw materials, and the degree of cell disintegration, must each be aligned with previous and subsequent processing steps in order to transfer the benefits of PEF application into enhanced mass transfer processes. In contrast to PEF pretreatments that cause cell disintegration, contact ultrasound is effective in reducing mass transfer barriers directly during drying. This leads to special requirements for equipment design and adaptations to the different drying processes, although recent advances have been made with contact ultrasound-assisted drying processes by using the product support for sound transmission. Screens can be used as efficient sound-transmitting surfaces when the ultrasound transducer is well matched with the screen frame. Such contact application enables ultrasound transfer in air drying and also in vacuum freeze- drying, where other mechanisms of sound transfer are greatly impaired. Further technical developments based on proven pilot concepts are necessary in order to implement a combined drying process on an industrial scale. The removal of water as one of the most important mass transfer processes in the food industry is also linked to a high consumption of energy, time, and costs. Depending on the temperature and pressure, the energy required to evaporate water is about 2.6 MJ kg 1. However, the total energy required for conventional drying is in the range of 4–6MJkg 1 of removed water, depending on the thermal efficiency of the drying system (which is in the range of 40–70%) and also on the dryer type, heat and mass transfer in the product, and heat losses (Singh, 1986; see Volume 4 of Modern Drying Technology). The modification of cell and tissue structures by PEF pretreatment or by simultaneous ultrasound application during drying demonstrated the potential to facilitate the removal of water, and resulted in increased diffusion rates and lower achievable moisture contents. Hence, the related reduction of drying time would result in energy savings and an improved utilization of production capacities, taking into consideration the moderate energy requirements for the PEF and ultrasound treatments. The synergetic effects of PEF pretreatment and ultrasound-assisted drying require further investigation. To reduce the barrier function of the cell membrane by PEF would result in a facilitated release of water from the intracellular space, whereas a subsequent drying process with the application of ultrasound may further improve the transfer of water and water vapor through the tissue. PEF pretreatments are effective at much lower energy inputs than are ultrasound pretreatments, and can ensure a uniform cell disintegration in all tissue layers. The disrupted cells might, in consequence, demonstrate an increased susceptibility towards the mechanical effects of ultrasound. Other combined treatments may include the application of microwave drying to PEF pretreated tissue, or in combination with ultrasound. Microwaves can overcome conventional limitations with regards to heat transfer due to volumetric heating. However, the improved heat transfer and the facilitated evaporation of water need to be accomplished by favorable mass transfer characteristics in order to benefit from the optimized 7.5 Conclusion 229 delivery of thermal energy to the product. The improved drying characteristics with a reduction in mass transfer barriers may allow the drying of larger pieces that are currently difficult to treat due to long drying times, inhomogeneous final moisture distributions, and quality losses under more severe drying conditions. Here, PEF pretreatment of the tissue or ultrasound application during drying may have further potential to contribute to process intensification.

Additional Notation Used in Chapter 7

A amplitude m c concentration kg m 3 E electric field strength V m 1 f frequency s 1 I intensity W m 2 MR moisture ratio – r cell radius m Rreflection coefficient – D Vm transmembrane potential V x travelled distance m Z acoustic impedance N s m 3

Greek Letters a damping coefficient – tan d ratio of elastic to viscous modulus – q orientation angle of cell in electric field rad

Subscripts and Superscripts rreflected 0 initial, incident, surface

Abbreviations

DM dry matter MRI magnetic resonance imaging PEF pulsed electric field us ultrasound 230 7 Process-Induced Minimization of Mass Transfer Barriers for Improved Drying

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8 Drying Assisted by Power Ultrasound Juan Andres Carcel, Jose Vicente García-Perez, Enrique Riera, Carmen Rossello, and Antonio Mulet

8.1 Introduction

In the current global context, the use of and demand for energy – and the short- term projections for both – highlight the need to include these matters in industrial process engineering. Moreover, environmental implications such as global warm- ing or air pollution have led to energy consumption becoming an issue that must be taken into account, not only by designing new unit operations but also by including modifications into the present systems. In this context, drying is among the oldest, most common and most diverse of all industrial operations (Mujumdar and Devahastin, 2000), with hot-air drying being one of the most widely used methods. In this technique, the moisture content reduction is achieved by changing the phase of water from liquid to vapor; this permits transport of the vapor from a solid or liquid/paste product to the surrounding gas medium, but implies a heavy energy demand (Lewicki and Michaluk, 2004). Moreover, as the thermal efficiency of the current industrial processes is low (25–50%), the amount of energy needed for the process is very high (Vega-Mercado et al., 2001) such that, today, drying operations utilize a very high proportion of the industrial energy production of developed countries (Mujumdar and Devahastin, 2000). The high temperatures used in hot-air drying may have negative effects on the quality aspects of some products, especially of colored or heat-sensitive components, and in these cases low-temperature drying methods should be considered to obtain high-quality products. A moisture reduction can be achieved by evaporation if the drying temperature is above freezing point, or by sublimation if the temperature used is below freezing point. Such low temperatures – especially those below freezing point – reduce the kinetics of any degradation reactions and maintain the natural attributes, thus achieving high-quality dehydrated products. This is certainly true in the case of vacuum freeze-drying (VFD), which is commonly known as freeze-drying, but also in the case of atmospheric freeze-drying (AFD). The low temperatures also reduce the drying kinetics, but in VFD this is partially overcome by applying vacuum conditions that increase the temperature gradient between the

Modern Drying Technology Volume 5: Process Intensification, First Edition. Edited by Evangelos Tsotsas and Arun S. Mujumdar. Ó 2014 Wiley-VCH Verlag GmbH & Co. KGaA. Published 2014 by Wiley-VCH Verlag GmbH & Co. KGaA. 238 8 Drying Assisted by Power Ultrasound

heating surface and the product, thus accelerating the drying process. Temperatures below freezing point allow heat-sensitive compounds to be preserved and also generate a highly porous structure in the final products, which show a great rehydration ability. Nevertheless, VFD has important rate-limiting steps linked to heat and mass transfer resistances. During the drying cycle, the product dried layer thickness increases, which in turn slows down the sublimation rate, whilst the long processing times require additional energy to run the compressor and refrigeration units. Another disadvantage is that vacuum conditions are easier to implement in batch than in continuous systems of production, thus impairing the economics of the operation (Kudra and Mujumdar, 2009). In AFD, water transport during drying is accelerated by circulating dry cold air aroundtheproduct,soastocreate a moisture gradient with a significant drying potential. The turbulence produced by the air circulation reduces the external resistance to mass transfer. Although this system allows continuous-mode operation, which in turn reduces the costs compared to VFD, the processing time is often very long, with significant energy needs for the dryer and peripheral facilities. Other drying techniques are based on decreasing the water load, not only by reducing the moisture content of the product but also by increasing the amount of solids. This occurs in the case of osmotic dehydration, where solid products are processed in hypertonic solutions of sugars or salts (Sutar and Gupta, 2007). The transport of water and solutes between the solid and the solution is also used to modify food characteristics, such as taste, color, nutritional value and preservation capacity. The rate of water loss and solute gain during osmotic drying depends on different variables, notably the solution concentration, temperature and physical characteristics, agitation of the osmotic solution, the structural characteristics of the solid, and the material–water interactions. In general, however, the process is characterized by the slowness of the solid–liquid mass transfer. As in the case of solid–gas drying techniques, the long processing times involve high facility, energy and labor costs. The intensification of drying must be mainly related to a shorter processing time. In hot-air drying, for example, increasing the air velocity will reduce the external resistance to mass transfer (Carcel et al., 2007a), whereas intermittent drying will reduce the energy consumption and provide products of a higher quality than those that are dried conventionally (Chua et al., 2003). The integration of different proven technologies – either new or hybrid – into existing plants might also contribute to drying process intensification. An infrared heat supply can, for instance, support hot-air drying and lead to shorter drying times by providing additional heat, so as to expedite the drying process (Chou and Chua, 2001). In microwave drying techniques, the applied microwaves produce internal heating and an increase of vapor pressure within the product that gently “pumps” the moisture to the surface (Turner and Jolly, 1991; Li et al., 2010; Lombrana~ et al., 2010), reducing the internal resistance to moisture transport. The use of heat- pump drying techniques in freeze-drying constitutes another example of process intensification with energy savings that can partially compensate the higher investment costs (Strommen and Kramer, 1994). Due to the slow nature of osmotic dehydration processes, several methods to speed up the kinetics have been 8.2 Ultrasound 239 considered (Rastogi et al., 2002). These include the application of pulsed-vacuum (Chiralt et al., 2001), pretreatments with electrical fields (Rastogi et al., 1999), or combinations of both techniques (Moreno et al., 2011). High-intensity ultrasound may represent an interesting alternative to intensify dehydration and dewatering in gas–solid and liquid–solid systems, affecting both internal and the external resistances to mass transfer (Mulet et al., 2003). The shorter processing time may result in a significant reduction of energy consump- tion; moreover, due to the fact that its effects are mainly mechanical, ultrasound does not induce any significant sample heating and, therefore, the product quality could be enhanced. There are many aspects of ultrasound to be considered, and many different approaches; consequently, it may be beneficial to summarize the effects found in different contexts to help develop further investigations of the subject. However, in order to achieve a better understanding of the intensifying effects on drying, it is important first to examine some of the characteristics of ultrasound.

8.2 Ultrasound

8.2.1 Ultrasound Waves

Acoustic energy is one of the types of energy that can be found in Nature. Sound consists of mechanical oscillations in time and space that take place in the material through which they travel. In this case, sound waves are elastic waves that need a material through which to be propagated (Carlin, 1960) but, unlike electromagnetic waves (e.g., microwaves), they cannot be transmitted in vacuum conditions. The frequency range of human hearing, from 20 Hz to 20 kHz, is used as the reference for the most common classifications of acoustic waves. Below this range the waves are termed as “infrasound,” and above the range as “ultrasound.” It should be noted that this classification is totally arbitrary and does not match any physical reasoning. Ultrasonic waves are characterized by the typical parameters of mechanical waves some of which, due to their importance in ultrasound applications, will be highlighted at this point.

8.2.1.1 Power Power is the total energy irradiated by an ultrasonic system per unit of time (J s 1 or W). Power is very important in ultrasound applications, due to the fact that the greater the power applied in a medium, the higher are the observed ultrasound effects. In general, it is convenient to relate the level of applied energy with the size of the considered system, thereby reporting the ultrasonic energy in a relative form. Thus, the ultrasonic intensity is defined by the power applied per unit of treated 2 area (W m ). The ultrasonic intensity, I, is related to the acoustic pressure, PA,by 240 8 Drying Assisted by Power Ultrasound

the relationship (Mason and Lorimer, 2002):

P2 I ¼ A ; ð8:1Þ 2rv where r is the density of the medium and v the wave velocity. Another useful way to express the power in relation to the size of the system is through the energy density (energy applied per unit volume, J m 3) that corresponds with the sound pressure level.

8.2.1.2 Frequency The frequency, f, is the number of cycles completed by a wave per unit of time. It is inversely related to the wavelength, l, which is the distance between two equally vibrating points:

¼ v : ð : Þ f l 8 2 Both, frequency and wavelength are of importance in some ultrasound effects. For example, the frequency is related to alternating compressions and expansions produced by ultrasound in a solid. Moreover, the relationship between wavelength and the pore size of the material defines their interaction; if the wavelength is greater than the diameter of a pore, this pore will be “transparent” to the ultrasonic energy, much like a boat being moved by a large wave in a storm. However, if the wavelength is in the range of the pore size, the sound wave can produce vibrations inside the pore, like small waves hitting the boat’s hull.

8.2.1.3 Attenuation During acoustic wave propagation through a medium, the wave intensity decreases as the distance to the emitter surface increases:

ad I ¼ I0e : ð8:3Þ a Here, I0 is the intensity emitted, d is the distance to the emitter surface, and is the attenuation (absorption) coefficient that is a characteristic property of the medium. Attenuation can be produced by wave reflection, dispersion or conversion into kinetic or thermal energy. An increase in frequency raises the attenuation coefficient (Mason and Lorimer, 2002).

8.2.1.4 Acoustic Impedance The specific acoustic impedance, Z, is the resistance of a medium to the propagation of a sound wave. It can be defined as the ratio of acoustic pressure to the so-called particle velocity at a single frequency (McClements, 1997). At an interface, the proportion of wave energy transmitted or reflected depends on the difference in impedance between the two media. Consequently, this difference determines the coupling between emitting surface and the treated medium. If the impedance difference is large, the proportion of energy reflected will be important and the ultrasound effects will be mainly localized at the interface. However, if the 8.2 Ultrasound 241 difference is small, a good coupling is achieved, a large portion of energy is transmitted, and the effects of ultrasound take place in the bulk of the medium.

8.2.2 Effects of Ultrasound on Mass Transfer

As pointed out, the characteristics of ultrasonic waves applied in a medium determine the effects produced. So, there are ultrasound applications that use low ultrasound intensities (<1Wcm 2) and high frequencies (>1 MHz) which do not affect the medium significantly. In these cases, the change in the characteristics of ultrasound waves traveling through the medium is used to determine the medium’s properties. As they constitute a nondestructive inspection technique, low-intensity ultrasound applications have been the subject of intensive research and development over the past few years (Benedito et al., 2002; Benedito et al., 2004; Ninoles~ et al., 2011). Another group of ultrasound applications considers the use of frequencies from 20 kHz to 1 MHz and intensities higher than 1 Wcm 2. The goal of these applications – known as power or high-intensity ultrasound – is to induce some effects in a process or product. In this sense, the application of ultrasound could be used to improve mass and heat transport and intensify processes such as osmotic or air drying. The effects produced by high-power ultrasound when applied in a medium are diverse, and their relative importance depends on the amount of energy transmitted and the characteristics of the medium itself. Ultrasound is pressure waves that produce alternating compression and decompression of the media. As these waves pass through a viscous medium, be it gas or liquid, it will dissipate this energy in the form of viscous flow (Riley, 2001). In liquids, when ultrasonic power attains a threshold, rarefaction may exceed the attractive forces between molecules and, from existing gas nuclei, cavitation bubbles can appear (Soria and Villamiel, 2010). These bubbles can maintain a stable cycle of increasing and decreasing size leading to so- called “stable cavitation” that generates a microagitation of the medium. However, the bubbles can also grow and collapse, giving rise to very high local temperatures (5000 K) and pressures (1000 atm) which produce, in turn, high-energy shear waves and turbulence in the cavitation zone. This last effect is known as “transient cavitation” (Leighton, 1998), and is observed at low frequencies (20–100 kHz). These dramatic microscale events can easily cause the breakdown of polymer chains or the damage of plant and animal cells (Kentish and Ashokkumar, 2011). Of particular relevance to many food applications is the occurrence of transient cavitation within the proximity of a solid surface. In this case, the bubble collapses asymmetrically and a microjet of fluid or bubbles (Lee et al.,2007)canbeemittedfromthebubbledirectlytowardthesolid surface, which can lead to pitting and erosion (Mason, 1998). This is the main effect observed in the use of high-intensity ultrasound in cleaning operations. Moreover, the microjets hitting the solid food surface may produce an injection of fluid inside the solid (Mason and Cordemans, 1996). The intensity of cavitation and its effects depend on the characteristics of the medium, such as viscosity, and/or process variables, such as ultrasonic intensity, ultrasonic frequency or pressure. 242 8 Drying Assisted by Power Ultrasound

The microjetting and the acoustic streaming patterns around solid objects are stronger within a few millimeters from the surface. These boundary layer effects are important, because it is usually such boundary layers that offer the greatest resistance to both heat and mass transfer. By concentrating the acoustic energy dissipation in these areas, ultrasound is extremely effective at improving heat and mass transfer kinetics, often more effective than other less site-specific options, such as high-shear mixers (Kentish and Ashokkumar, 2011). In gas media, the main challenge of ultrasound application is to attain an efficient transmission of the acoustic waves at high frequencies, due both to the acoustic impedancemismatchbetweentransducersandgasaswellastothehighultrasonic attenuation in gas media. This explains why, at present, only a few research groups are working in this field and there are few authors who have published results that show successful applications. However, when ultrasound is applied in an efficient way, it can produce intense effects on the interfaces, such as pressure variations, interfacial instabilities or microstirring, which can also affect the mass and heat transfer phenomena (Carcel et al., 2007a) by reducing the boundary layer thickness. In solid materials, alternating compressions and expansions generated by the ultrasonic waves induce a similar effect to that observed when a sponge is squeezed and released repeatedly (De la Fuente et al.,2006).This“sponge effect” helps to release liquid from the inner part of a particle to its surface and promotes the suction of fluid from outside. The forces involved in this mechanism can be higher than the surface tension which maintains the water molecules inside the capillaries of the material, creating microscopic channels (Muralidhara et al., 1985) and facilitating the exchange of matter. Other effects to be considered are the variation of viscosity, surface tension or the deformation/degradation of the solid structure. From a general point of view, all of the effects produced by ultrasound could influence mass and/or heat transfer phenomena. In the treatment of a solid immersed in a fluid, ultrasound can accelerate the internal transport, making it easier for fluids to enter and/or exit the solid matrix, and it can also promote the exchange of matter between the solid surface and the surrounding fluid. Consequently, the use of ultrasound, when applied in an efficient manner, could be interesting in applications involving heat or mass transport, decreasing both the external and internal transport resistances. These effects could be used to intensify operations such as drying. Moreover, ultrasound is considered by the general public as safe, nontoxic and environmentally friendly which, in areas like the food industry, represents an important advantage over other processing techniques, such as microwaves, pulsed electric fields or gamma radiation, which people may view cautiously (Kentish and Ashokkumar, 2011).

8.3 Ultrasonic Equipment

Any system to apply ultrasound must be able to produce a stable and reliable ultrasonic field. Usually, it is composed of three basic elements: a source of energy; 8.3 Ultrasonic Equipment 243 a transducer system; and a system for the transport and application of the ultrasonic waves.

8.3.1 Source of Energy

At first, any type of energy is susceptible to be converted into ultrasonic energy; for example, the mechanical energy accumulated in fluids could be transformed into ultrasonic waves by means of a siren. However, the main applications at the industrial level include systems that use electricity as the energy source.

8.3.2 Transducers

The transducer system is made up of the devices used to convert the energy provided into ultrasonic energy. There are different types of transducer, the characteristics of which depend on the source of energy available, the medium where ultrasound must be applied, and the specific application considered. In order to transform the kinetic energy of a fluid into ultrasonic waves, siren and whistle systems are used. In siren systems, the fluid is forced to pass across a hole and depending on its velocity, the fluid flux will generate turbulence that constitutes the mechanical wave. In whistles, the fluid is forced across a thin blade which causes the blade to vibrate. For each vibrational movement, the leading face of the blade creates a pressure wave (Mason, 1998). Siren and whistle systems are not widely used in industrial applications. In liquid applications, the whistles constitute a powerful tool for mixing and homogeniza- tion (Mason and Lorimer, 2002), and for this reason they have been applied to cheese-making processes or in the sauce industry to achieve stable emulsions, such as mayonnaise or ketchup (Mason, 1998). Sirens are very rarely applied in a gas medium, most likely due to the fact that the working frequency of these systems is within the human hearing range, so that noise makes the industrial operation difficult. An example was provided by Da Mota and Palau (1999), who used a siren system to improve onion drying at low frequencies (1.6 and 3.2 kHz). If the energy source is electrical, two main types of transducer are available: magnetostrictive and piezoelectric. The magnetostrictive transducer is based on the changes of shape produced in some high-strength metallic alloys by a changing magnetic field. This material shows a reduction in dimension when applying a current (magnetostriction) and, consequently, a reduction in the dimensions of the transducer. Switching off the electrical current leads to a return to the original dimensions, and the repeated rapid switching on and off of the current generates the mechanical vibrations required (Mason, 1998). The dimensions of the transducer must be accurately designed so that the whole unit resonates at a precise sound frequency. The piezoelectric transducer is based on the piezoelectric effect presented by some materials. This is in reality composed of two contrary effects, the direct and the 244 8 Drying Assisted by Power Ultrasound

Fig. 8.1 Piezoelectric effect. (a) Direct; (b) Reverse.

reverse effect (Fig. 8.1). The direct effect appears when pressure is exerted on the material, creating an equal electrical charge in each face of the material but with opposite signs. The reverse effect occurs when an electrical charge, opposite in sign, is applied to the two faces of the material. The different signs of the electrical charges make the material contract (or expand, depending on polarity), creating a mechanical wave. If the polarity changes the material expands (or contracts), creating a new mechanical wave. These changes in shape, at frequencies higher than 20 kHz, generate an ultrasonic vibration. The piezoelectric materials used in the transducers are usually ceramic, made from randomly oriented ferroelectric crystals and metallic elements for mechanical and overheating protection (Gallego-Juarez, 1998). In high- power ultrasound applications, one of the most commonly used arrangements is the sandwich type, which consists of two (or more) ceramic pieces, connected in opposite directions of polarity, between two metallic cylinders. The structure is compressed by an axial bolt, and the cylinder length is linked to the signal wavelength. Magnetostrictive and piezoelectric transducers have both advantages and draw- backs. Magnetostrictive transducers are able to reach high levels of acoustic power intensity, over 150 Wcm 2, the acoustic field produced is very stable and reliable, and they do not age (Peshkovsky and Peshovsky, 2010). A drawback is that they present a relatively low energy efficiency (<50%), which is especially important when compared to piezoelectric systems that achieve rates of conversion of up to 95%. Moreover, the energy losses produce transducer heating, making it necessary to consider external cooling and, in addition, restrict the frequency range to below 100 kHz. In contrast, as noted above, the electrical–ultrasonic efficiency of the piezolectric transducer allows energy to be saved when compared to magnetostrictive transducers. This is probably the main reason why the former is more widely used than the latter, regardless of the relatively low levels of acoustic power intensity provided and the short life-span. Additionally, piezoelectric transducers are high- voltage-driven and air-cooled which, for some applications, implies the risk of explosion (Peshkovsky and Peshovsky, 2010). It is important to emphasize that the great majority of the existing high-power ultrasonic transducers are still based on the classical piezoelectric sandwich transducers which were originally designed by Langevin in 1920. These transdu- cers have many limitations to meet industrial requirements (Gallego-Juarez et al., 2010), but some of these problems have been resolved by appropriate designs of the application system. 8.3 Ultrasonic Equipment 245

8.3.3 Application Systems

Usually, the transducers are attached to a vibrating system, the function of which is to transmit the vibration from the transducer to the medium. The design of this part of the ultrasonic system is fundamental for optimal transmission of the ultrasonic energy to the medium. From the strict point of view of design, the transducer and the vibrating system must be jointly considered, due to their close interaction. Nevertheless, in this chapter, they are considered separately so as to better illustrate the operation of these ultrasonic devices. It is important to address this point because, on many occasions, the use of an ultrasound effect is discarded because of energy transmission problems. One of the main problems of ultrasound applications is to achieve a good coupling between the generation system and the medium to be treated. The impedance difference between the vibrating system and medium makes the transmission of ultrasonic waves difficult and causes wave reflection that could even damage the generation system. Hence, the greater the impedance difference, the lower is the efficiency of ultrasound transmission. An optimal design of the application systems could achieve a better coupling, thus reducing the reflections produced by the impedance difference. The next problem for the proportion of wave transmitted to the medium is the attenuation of the signal by energy absorption. This is a characteristic of each specific medium, but the design of the vibrating system could also help to reduce its effect.

8.3.3.1 Treatments in Liquid Media In liquid applications, coupling is relatively easy due to the fact that the impedance difference between the vibrating systems (solid) and the medium (liquid), although important, does not represent such a major problem as in the case of gas applications. To the most widely used vibrating systems belong the so-called probe- type systems or horn systems, where the ultrasonic vibration is directly applied by a vibrating “horn,”“probe,” or “sonotrode.” Depending on the geometry of the probe (Fig. 8.2a and b), it could be used simply to transmit the ultrasonic energy (cylindrical) or to concentrate it on a smaller surface (conical, exponential or stepped) in order to amplify the intensity and, therefore, the ultrasonic effects (Mason, 1998). The final size of the probe tip depends on the gain factor of the probe, and reduces as the gain factor increases. These arrangements allow levels of high ultrasonic intensity to be obtained that generate important ultrasound effects. However, the energy concentration on the small surface of these tips produces violent cavitation that could erode the tip and limit its operating life. Another important drawback of these systems appears when applied on an industrial scale where large volumes must be treated, as the cross-sectional area of the horn tip in contact with the load is small. This problem may be overcome by using arrays of sonotrodes. Probe systems have been used to study the influence of ultrasound application on osmotic dehydration, for example the treatment of apples in sugar solutions (Carcel 246 8 Drying Assisted by Power Ultrasound

Fig. 8.2 Ultrasonic systems for liquid applications. (a) Sound probe system; (b) Sonotrodes different in shape and emitter surface; (c) Semi-industrial ultrasonic bath.

et al., 2007b), or meat brining (Carcel et al., 2007c, Siro et al., 2009). They have also been used to apply ultrasound as pretreatment in liquid media prior to air-drying. For example, Rawson et al. (2011) used a 1.5 kW ultrasonic processor at 20 kHz with a 19 mm diameter probe to apply ultrasound in distilled water as a pretreatment for carrot followed by hot-air drying or freeze-drying. Another common application system in liquid media is the ultrasonic bath (Fig. 8.2c). Currently, commercial ultrasonic baths are available for laboratory or industrial applications, with different sizes, ultrasonic power or frequency. All of this equipment consists of an array of transducers that are attached to a plate and which vibrate in phase; this, in turn, makes the plate vibrate. In some industrial applications, the plate is placed at the bottom of a tank where the treatment takes place. Typically, the transducer–plate set is isolated from the content of the tank, but this is an adaptable arrangement and by placing a larger 8.3 Ultrasonic Equipment 247 number of transducers on a plate with a larger surface area a greater amount of liquid can be treated. In commercial ultrasonic laboratory baths, the transducers are directly attached to the metallic tank that constitutes the bath. Then, when the transducers vibrate, not only does the bottom of the tank vibrate but also the bath wall, creating a chaotic ultrasonic field in the liquid inside. Although the waves are generated in different directions, a stationary ultrasonic field with regions of maximum and minimum ultrasonic energy is created. As the ultrasonic field is not uniform, the effectiveness of the ultrasonic treatment depends on the location of the samples in the bath. Baths have been used to apply ultrasound during the osmotic dehydration of apples in sugar solutions (Simal et al., 1998, 2006), cheese (Sanchez et al., 1999) or meat brining (Simal et al., 2006), and also to study its effects on mass transport kinetics. Ultrasound has also been applied in liquid–solid systems as a pretreat- ment prior to osmotic dehydration or hot-air drying of products such as banana (Fernandes and Rodrigues, 2007), pineapple (Fernandes et al., 2008) or malay apple (Oliveira et al., 2011).

8.3.3.2 Treatments in Gas Media The application of high-intensity ultrasound in gas media with the aim of intensifying the drying is more difficult than in liquid media, due to gases being highly attenuating media that absorb the acoustic energy, preventing its transfer to the solids being treated. Additionally, the high impedance difference between the solid surface of emitters and the gas, and between the gas and the solid samples, produces the reflection of a high proportion of the generated acoustic signal (García-Perez et al., 2006a). Therefore, the main drawback of these treatments is transmission of the acoustic wave from the emitter’s surface to the samples. Bantle and Eikevik (2011) used a commercial vibrating transducer to apply ultrasound during the air freeze-drying of peas, and found the drying time to be shortened by about 10%. This small reduction not only showed the positive effect that ultrasound application has on drying kinetics, but also highlighted the difficulty of generating and transmitting a high-intensity ultrasonic field in gas media using existing commercial ultrasonic systems. Another approach to ultrasound-assisted convective drying was reported by Schossler€ et al. (2012), who used a ring sonotrode driven by an ultrasound processor originally designed for ultrasonic sieving. The ring sonotrode was equipped with a laboratory drying screen that served as the sample-supporting and sound-transmitting surface. The whole system was placed in a drying oven. With this system, it was found that the drying time needed to reach a residual moisture content of 20% for red bell pepper and apple cubes was reduced by 18% and 27%, respectively. In this arrangement, problems of the transmission of ultrasound through the air were partially resolved by a direct contact of the vibrating screen and the sample. Hence, strictly speaking, it is not an application of ultrasound in a gas–solid system but rather an application in a solid–solid system, namely the application system and the sample. In this sense, the ultrasonic vibration applied in direct 248 8 Drying Assisted by Power Ultrasound

contact between the emitter surface and sample has obtained good results (Gallego-Juarez, 1998). Contact between the vibrating system and the food material leads to a good acoustic impedance matching that favors a deep penetration of the acoustic energy and increases the effectiveness of the process. Gallego-Juarez et al. (2007) developed a prototype that also applies static pressure to obtain a good homogeneous mechanical coupling between the sample, the vibrating tip of the transducer, and a porous layer that constitutes the sample holder. From the modeling of the drying experiments, these authors identified diffusivities for apple ranging from 2.93 10 8 m2 s 1 withoutultrasoundupto8.84 10 7 m2 s 1 with ultrasound application (50 W). Values for potatoes ranged from 2.2 10 8 m2 s 1 without ultrasound up to 1.37 10 7 m2 s 1 with ultrasound, which implied an increase of one order of magnitude. However, the direct contact between emitter and sample makes it difficult to implement this system on an industrial scale. It is of great importance to gain further insight into the mechanisms of ultrasonic wave transmission in gas media in order to optimize the application systems (De la Fuente et al., 2006). In this sense, a new family of power transducers with extensive radiators has been developed (Gallego-Juarez, 1998; Gallego-Juarez et al., 2001). The extensive surface area of these transducers increases the radiation resistance and offers the vibration system a good impedance matching with the gas medium (Fig. 8.3). These radiators are generally circular or rectangular vibrating plates with a stepped, grooved or stepped-grooved profile, although they might also be of any shape adapted to the specific problem, such as a cylinder, a shell or any other (Gallego-Juarez et al., 2010). The special shape and profile of the radiator is designed to control the vibration distribution and the radiation pattern, so as to achieve a large amplitude of vibration and highly directional or focused beams for energy concentration. Moreover, the high power capacity and the extensive radiating area make it easy to apply at industrial scale. These systems have been used to apply ultrasound in diverse operations, such as defoaming in reactors or canning lines, or in the agglomeration of particles for air decontamination. The application of airborne ultrasound during drying has been carried out with circular and rectangular plates (Gallego-Juarez et al., 2010), but mainly with cylindrical systems (Riera et al., 2011). The latter system has been tested during the convective drying of several fruits and vegetables under different process conditions, such as air temperature (García-Perez et al., 2006b) and velocity (Carcel et al., 2007a), mass load (Carcel et al., 2011b) or applied ultrasonic power (García-Perez et al., 2009), achieving reductions in the drying time of between 30% and 75% and increases in diffusivity of up to 400% depending on the product. These results showed that the efficient application of ultrasound in convective drying could be an interesting way to intensify the drying operation. However, attention must be paid not only to application systems but also to the interaction between ultrasonic waves and drying chambers. The correct design of drying walls and the use of ultrasonic wave reflectors could guide, or even 8.3 Ultrasonic Equipment 249

Fig. 8.3 Ultrasonic systems for application in gas media. (a) Cylinder; (b) Stepped plate.

concentrate, the ultrasonic energy towards the samples, thereby increasing their effects. Some attempts have been made to include these aspects in the application systems by Gallego-Juarez et al. (2010) and Khmelev et al. (2011). Thus, it can be concluded that the correct design of an ultrasonic system (generator, transducer and application systems) represents a key step towards the efficient application of ultrasound, and generates significant effects in drying. However, it is also extremely important to appreciate the influence that different process variables have on the levels of drying intensification linked to ultrasound application; these aspects are addressed in the next section. 250 8 Drying Assisted by Power Ultrasound

8.4 Influence of the Main Process Variables on Drying Intensification by Ultrasound

As noted above, any technological process – specifically drying – may be subject to process intensification, and high-intensity ultrasound application could be an interesting alternative by which to perform this operation. However, the different process variables may influence exactly how extensive are these effects produced by ultrasound and, in turn, affect aspects such as drying kinetics, energy consumption or final product quality. The first group of variables involves those related to high- intensity ultrasound application. As the intensifying effects of ultrasound depend on the characteristics of the applied wave, consideration must be given to properties such as the actual applied ultrasonic power, the attenuation that the wave experiences when passing through the treated medium, or the impedance difference between the application system and the media. A second set of variables to be considered is related to conventional drying processes, such as temperature or the level of turbulence of the drying agent. Usually, the higher the temperature, the faster is the drying kinetics and the greater can be the damage to thermosensitive compounds. The increase in turbulence level reduces the external resistance and increases the mass transfer rate though, in air drying, it also favorstheappearanceofcase-hardeninginsomeproducts.Thesevariablesaffectthe level of energy in the medium and might also influence the characteristics of the ultrasonic wave and its capacity to induce effects on mass transport. The final group of variables to be considered is related to the inlet and outlet process streams, and their respective properties (enthalpy, moisture content, temperature, viscosity, thermal conductivity, etc.). In drying, the inlet streams are mainly the drying material and the drying agent (air in convective drying, hypertonic solution in osmotic drying, etc.), while the outlet streams are the product and the outlet drying agent. As far as the properties of the feed are concerned, some are characteristic of the material to be treated and, thus, are not prone to change to intensify the process (e.g., porosity, viscosity, or thermal conductivity). However, these properties can exert an influence on both the drying kinetics and final product quality, and may make the intensification action of ultrasound either easier or more difficult. For example, the structure of the material could condition the effects of ultrasound. Clearly, a large number of variables – including product and process characteristics – are involved when ultrasound is applied withtheaimofintensifyingthedrying, and hence it is difficult to draw general conclusions. For these reasons, the influence of the most important variables is addressed in the following subsections, and illustrated with examples of applications.

8.4.1 Ultrasonic Power Applied

The ultrasound wave characteristics influence the observed effects of ultrasound when applied to a medium. The wave frequency is linked to the microstirring 8.4 Influence of the Main Process Variables on Drying Intensification by Ultrasound 251 produced by ultrasound at the interfaces, the effects of which become more marked as the applied frequency rises. However, another wave characteristic – the attenuation – increases along with the frequency (Mason and Lorimer, 2002), reducing the amount of available energy. It is for this reason that the frequency used for high-intensity ultrasound applications is relatively low (20–100 kHz). The attenuation produced by the drying agent also affects the available energy, in the same way as the impedance difference between application systems and the treated media. Therefore, the main factor responsible for ultrasound effects is the actual ultrasonic energy received by the sample. As a consequence, it is interesting to establish the relationship between the actual ultrasonic power applied and the intensification effects induced by ultrasound.

8.4.1.1 Ultrasonic Field Measurements Due to the fact that the effects of ultrasound are directly related to the ultrasonic field applied, its characterization is fundamental; only then is it possible to evaluate the influence of ultrasound application and to establish clear relationships between ultrasonic parameters and observed effects. Nowadays, it is still common for ultrasonic power to be reported as either the electrical consumption of the generator or as the electrical energy supplied by the generator to the transducer. However, this information is incomplete because the ultrasonic field actually applied to the product also depends on other parameters, such as the electric– ultrasonic conversion efficacy of the transducers (Lin and Zhang, 2000), the size and geometry of the applying chamber, or the physical properties of the treated medium (e.g., viscosity, surface tension, vapor pressure, composition, temperature, impedance matching) that define its wave attenuation capacity (Mason and Lorimer, 2002). Therefore, the amount of ultrasonic energy received by the sample cannot be exclusively expressed as the electrical energy consumed by the system and, from that, the need arises to use adequate methods to characterize the actual ultrasonic field applied in each specific process. The parameters that best characterize an ultrasonic field are the wave frequency and the ultrasonic energy applied, with the latter being expressed as power, intensity, or acoustic pressure. The frequency does not represent a critical point because the variation of the applied frequency is usually less than 5–10% of the nominal frequency provided by the manufacturer. In contrast, the actual applied acoustic energy could be very different from the electrical consumption, and its direct measurement is difficult. In the literature, it is possible to identify a wide variety of methods for that purpose, but the results provided may be difficult to compare one with another. The majority of these methods are based on the measurement of physical or chemical changes produced by ultrasound, and a classification was proposed by Berlan and Mason (1996):

Calorimetric methods. Electrical and/or mechanical measurement of transducers. 252 8 Drying Assisted by Power Ultrasound

Methods based on direct mechanical effects, including measurements using acoustic probes, optical methods and acoustic impedance, radiation forces, liquid surface distortion or particle velocity measurements. Methods based on secondary effects produced by the propagation of sound and cavitation. This group includes methods based on volume change, noise measurements, conductance changes, electrical effects, sonoluminescence and chemical probes.

In the case of liquid applications, the calorimetric methods should be emphasized. A simple calorimetric method consists of measuring the tempera- ture increase during the first few moments (90 s) of ultrasound application (Raso et al., 1999). The procedure assumes that, in this early stage, the system behaves in a quasi-adiabatic manner, and the temperature variation is produced only by the application of ultrasound. The applied ultrasonic power, P, can then be calculated from: dT P ¼ Mc ; ð8:4Þ dt where M is the liquid mass, c is its heat capacity, and dT/dt is the temperature change. This method provides a measurement of the ultrasound energy introduced into the medium and converted to heat. This is not a direct measurement of the applied ultrasonic field, but it could constitute a reference for comparison with the results of other methods, and also be used to quantify the effect of ultrasound. Another popular method for measuring the ultrasonic field is the use of piezoelectric direct effect-based systems, such as microphones (García-Perez et al., 2006a) in gas applications or hydrophones (Carcel et al., 2007c) in liquid applications, due to their simplicity and reliability. The microphones/hydrophones are electroacoustic sensors that allow the measurement of not only the acoustic pressure but also other acoustic field parameters, such as frequency, without disturbing the field (Kutruff, 1991). The figures obtained are local values that permit the detection and characterization of stationary fields and the location of maximum and minimum ultrasonic intensity zones (Pugin and Turner, 1990). Occasionally, the signal received must be filtered to avoid the noise masking the measurement. In liquid media, intense cavitation could damage the sensor. Moreover, online measurements of the electrical (phase, intensity, power, voltage or frequency) and/or the mechanical (vibration amplitude) response of transducers during the ultrasound application could also provide information about the correct operation of the system (Gallego-Juarez et al., 2007) and permit the detection of system failures.

8.4.1.2 Ultrasonic Intensity and Effects It is known that the applied ultrasonic energy determines the extent of ultrasound effects. As stated before, a widely accepted classification of the industrial ultrasound applications divides these into two groups, low-intensity and high- intensity applications. While in the former case the ultrasonic intensity used is 8.4 Influence of the Main Process Variables on Drying Intensification by Ultrasound 253

under 1 Wcm 2 to avoid significant changes in the products, the latter case uses higher intensities to induce effects in the treated materials. The effects may differ when ultrasound is applied through a liquid or gas medium.

Liquid–Solid Systems In liquid–solid systems, such as solids treated in hypertonic solutions, two main forms of mass transport can be identified: (i) water moving from the solid tissue to the solution; and (ii) solutes moving from the solution into the solid. Both fluxes must overcome the internal resistance to mass transfer in the interior of the treated solids, and the external resistance between the solid surface and the surrounding liquid. The application of high-intensity ultrasound in these systems can affect moisture and solute transport, reducing not only the external resistance to mass transfer but also the internal resistance. The influence on internal resistance comes from the appearance of ultrasonic effects, such as the above-mentioned “sponge effect” or the creation of microchannels, which can make the movement of solutions inside the solid easier. At the outer solid–liquid interface, the generation of intense microstirring or the implosion of cavitation bubbles affects the external resistance by reducing the boundary layer of diffusion (Carcel et al., 2012). As a consequence, the samples undergo a faster osmotic dehydration, increasing the solid content and reducing the moisture load. All of these effects are intensity-dependent. If the applied ultrasonic intensity is low, the energy applied may be insufficient to generate effects with a significant impact on mass transport. Even the appearance of some of these effects, such as cavitation or the generation of microchannels, requires a minimum level of supplied energy, and therefore an ultrasonic intensity threshold exists that must be applied to produce significant effects to affect the mass transport. Above this threshold, the increase in ultrasonic intensity induces more intense effects, which have a greater influence on the mass transport kinetics. This fact has been observed in different solid–liquid systems. Thus, when studying the ultrasonically assisted osmotic dehydration of apple in a 30Brix sucrose solution, Carcel et al. (2007b) observed that both main mass fluxes, water and solids, were affected by the application of high-intensity ultrasound. The extent of the effects of ultrasound depended on the applied ultrasonic intensity. In this way, for the same treatment time (45 min), the higher the ultrasonic intensity applied, the lower was the water content and the greater the increase in dry matter in the samples (Fig. 8.4). However, below an ultrasonic intensity of 8 Wcm 2 the changes in the moisture and solid contents of the samples were not significant, showing that a minimum ultrasonic intensity thresholdmustbeappliedtoobservesignificant effects. At the maximum ultrasonic intensity tested (11.5 Wcm 2), the water effective diffusivity was 117% higher than that identified for conventional osmotic drying experiments. In the case of solid transport, ultrasound produced an effective diffusivity increase of 137%. Moreover, the increase in dry matter was significantly higher than that observed in experiments carried out by applying an intense agitation of the osmotic solution. This fact shows that the intensification of mass transfer produced by ultrasound is greater than that produced by conventional methods, 254 8 Drying Assisted by Power Ultrasound

-0.6 0.7

-0.8 0.6

-1.0

0.5

-1.2 Increase in dry matter Increase (kg d.m./kg initial d.m.)

(kg water/kg initial d.m.) 0.4

Decrease of moisture content -1.4

-1.6 0.3 0 4 8 12 Ultrasonic intensity (W cm-2)

Fig. 8.4 Decrease in moisture content and increase in dry matter of apple samples treated for 45 min in a 30 Brix sucrose solution with ultrasound application at different intensities.

such as the increase in osmotic solution turbulence by stirring to reduce the thickness of the boundary layer for diffusion. The applied ultrasonic intensity may also influence the relative importance of the different effects generated by ultrasound. Thus, when high ultrasonic intensity levels are applied, effects such as interface microstirring are enhanced. However, the implosion of cavitation bubbles near the solid becomes also more violent, generating powerful microjets of solution that impact the solid surface (Mason and Lorimer, 2002). The more powerful effects produced by cavitation might mask the effects produced by microstirring. The cavitation microjets could also result in the microinjection of osmotic solution inside the solid, increasing the solute content significantly and, as a consequence, the solute transport kinetics would be significantly increased by the application of ultrasound. However, the injection of solution makes it difficult for the internal moisture to flow out, decreasing the water transport kinetics; in fact, if sufficient ultrasonic intensity is applied the product’s moisture content may even be increased at the end of the treatment. When Siro et al. (2009) used a probe system to study ultrasound application during pork loin brining, ultrasonic treatments in the range of 2–4Wcm 2 were found not to have any significant effect on the water-holding capacity of samples with respect to statically brined samples. In contrast, the salt transport was dependent on the ultrasonic intensity applied, as the identified diffusion coefficient of NaCl rose as more ultrasonic intensity was applied in the mentioned range of 2–4Wcm 2. Ozuna et al. (2010) also studied the pork loin brining by using an 8.4 Influence of the Main Process Variables on Drying Intensification by Ultrasound 255

3.8 50 kg m-3 ; without ultrasound 50 kg m-3 ; with ultrasound 3.6

3.4 O /kg initial d.m.) 2 3.2 Initial moisture content of meat

3.0

2.8 280 kg m-3 ; without ultrasound Moisture content (kg H 280 kg m-3 ; with ultrasound 2.6 0 907560453015 135120105 Time (min)

Fig. 8.5 Evolution of the moisture content of pork loin samples during treatment in brines of different NaCl concentrations, with and without ultrasound application.

ultrasonic bath (150 W nominal power, volume 6 l) and brines with different NaCl concentrations. It should be noted that the meat moisture content decreased with time for experiments carried out with a brine salt concentration of 280 kg m 3, whereas it was increased at 50 kg m 3, because the brine salt concentration influenced the water-holding capacity of meat (Siro et al., 2009). Ozuna et al. (2010) found that changes in the experimental moisture content of the meat with time, with and without ultrasound application, were not significantly different (see Fig. 8.5). From modeling, the differences between the effective diffusivity identified (from 1.04 10 10 m2 s 1 without ultrasound to 2.60 10 10 m2 s 1 with ultra- sound) were not significantly different (confidence level >95%), but this can be mainly attributed to experimental scatter and natural raw material variability. Nonetheless, this illustrates how important it is not only to carry out the experiments carefully but also to perform a large number of replicates to draw statistically significant conclusions. However, as regards the NaCl transport, the experimental variability was smaller and an increase by 45% on the identified NaCl effective diffusivity when ultrasound was applied was found to be significant. The ultrasonic intensity applied by both Siro et al. (2009) and Ozuna et al. (2010) was relatively low. Carcel et al. (2007c) used higher intensities during meat brining for 45 min in a saturated NaCl solution, and found a reduction in water losses when the ultrasonic intensity level was above 51 Wcm 2 compared to convention- ally brined samples. Moreover, the samples treated at the highest acoustic intensity tested (75.8 Wcm 2) showed a higher moisture content than that of fresh meat, 256 8 Drying Assisted by Power Ultrasound

which indicated that these samples were not dehydrated but hydrated. With regards to the NaCl content, this was 115% higher in ultrasonically brined samples, and therefore at high ultrasonic intensity levels the application of ultrasound had induced both water and NaCl gains. The highest applied acoustic intensity generated a more intense cavitation that could have caused the microinjection of brine into the meat samples. This fact would explain not only the increase in NaCl content but also the rise in sample water content by the increase of brine content inside the meat sample. The magnitude of the cavitation effect most likely also masked other effects produced by ultrasound, affecting the net mass transport of moisture and NaCl.

Gas–Solid Systems In gas–solid systems, such as air-drying processes, the effective application of ultrasound can contribute to significantly shorter processing times. This effect has also been reported for different products such as carrots (García-Perez et al., 2008), lemon peel (García-Perez et al., 2009), olive leaves (Carcel et al., 2010), potatoes (Ozuna et al., 2011a), or eggplant (García- Perez et al., 2011). Effects induced by the application of ultrasound, although causing changes to both external and internal mass transport resistances, are somewhat different than in liquid–solid systems. Notably, no cavitation takes place on gas–solid interfaces; rather, if cavitation appears it will be in the liquid phase inside the solid. The reduction of internal resistance is mainly associated with mechanical effects brought about by applying ultrasound to the material being dried. The alternating cycles of expansion and contraction (“sponge effect”) may help to make it easier for water to flow out of the solid matrix. Moreover, the mechanical stress could lead to the creation of microchannels, which eases water/vapor movement inside the solids. As far as the influence on external resistance is concerned, it could be linked to a reduction of the boundary layer thickness due to different effects, such as pressure variations, oscillating velocities and micro-streaming on the solid–gas interfaces. These effects should be responsible for reductions of the boundary layer of diffusion and, as a consequence, for improvements in the vapor transfer rate from the solid surface to the air medium. How widely extended are the effects produced by ultrasound also depends on the power applied, with a minimum level of ultrasonic power needing to be provided if the drying kinetics is to be significantlyaffected.Thus,García-Perez et al. (2009) studied the drying of carrot and lemon peel at 40 Cand1ms 1 by applying ultrasound at different power densities (variation in nine steps from 0to37kWm3). Here, the kinetics of drying was seen to be intensified by ultrasound application, and the ultrasound effects depended on the power applied. For carrots, the drying time needed to reach an average moisture content of 1.5 kg water per kg dry matter (average initial moisture content 8.26 0.45) was reduced by 32% when the applied acoustic power density was 37 kW m 3. However, the influence of ultrasound on experimental drying kinetics was not significant until the applied acoustic power density exceeded a minimum value of between 8 and 12 kW m 3. 8.4 Influence of the Main Process Variables on Drying Intensification by Ultrasound 257

The influence of applied ultrasonic power on drying kinetics could be quantified by modeling the experimental data. To this purpose, a diffusion model which considers external resistance to be negligible provides a relatively simple model that could be very useful as a first approach to modeling the drying process. Hence, the identified moisture effective diffusivity includes every aspect that has an influence on kinetics. By using this model, García-Perez et al. (2009) observed that the greater the acoustic power density applied, the higher were the values of effective diffusivity obtained during carrot drying. Moreover, García-Perez et al. (2009) found that, above the ultrasonic power density threshold, the values of effective moisture diffusivity identified for carrot drying were proportional to the applied acoustic power density, and this led to the establishment of a significant linear relationship between both variables. A similar behavior was found when lemon peel was dried under the same experimental drying conditions. Unfortu- nately, the fit of this type of model to experimental data might not be adequate when the assumed boundary conditions do not correspond to the actual experimental conditions. Thus, in the particular case of studies performed by García-Perez et al. (2009), the low value of explained variance obtained (i.e., the percentage of data variability explained by the model) of about 90% could be attributed to the fact that the air-drying velocity used in the experiments (1 m s 1) made it difficult to consider the external resistance to mass transfer as negligible. Consequently, this model would permit only a rough quantification of the influence of power ultrasound on drying kinetics. Another important drawback was that the simplified model included all of the factors affecting the drying rate in the effective diffusivity identified; hence, it was not possible to distinguish between the influence of ultrasound on the external and the internal resistances. By changing the outer boundary condition, external resistance can be included in the diffusion model, though this makes the mathematical resolution of the model more difficult. In this case, two parameters must be identified: (i) the effective diffusivity (internal resistance); and (ii) the mass transfer coefficient (external resistance). In order to reduce the number of possible solutions, the effective diffusivity identified by neglecting the external resistance can be used as the start value of diffusivity. The diffusion models that include external resistance fit the experimental data better and, therefore, provide a better description. Thus, García-Perez et al. (2011) fitted both types of model to the experimental data of drying kinetics of eggplant (40 C and 1 m s 1) at different ultrasonic power intensities. The percentage of explained variance was improved from an average value of 85% to 99.9% when the external resistance was included in the model. Moreover, these models also enable the separate quantification of the influence of ultrasound application on the internal and the external resistances to mass transport. In this way, García-Perez et al. (2011) found that both the effective diffusivity and the mass transfer coefficient identified for eggplant drying were affected by ultrasound application. Moreover, these effects were more marked the higher the acoustic power density applied. Thus, the effective diffusivity identified for conventional drying was 8.9 1.3 10 10 m2 s 1, whereas when an acoustic power density of 37 kW m 3 258 8 Drying Assisted by Power Ultrasound

was applied the identified diffusivity was 27.9 3.0 10 10 m2 s 1, which implies an increase of 220%. With regards to the mass transfer coefficient, the observed increase was similar when ultrasound was tested at the maximum power density (37 kW m 3), from 1.87 0.4 10 3 kg water m2 s 1 in conventional drying to 6.16 0.9 10 3 kg water m2 s 1. Among the few reports that have focused on the influence of applied power on drying kinetics parameters, it has been stated that, above the ultrasonic power threshold, the relationship between the level of applied power and effective diffusivity or mass transfer coefficient is linear. In these studies, the applied power ranged from 0 to 37 kW m 3. García-Perez et al. (2011) identified significant (confidence level >95%) linear relationships between diffusivity or mass transfer coefficient and the ultrasonic power density applied, in both cases achieving correlation coefficients of over 0.98. Such a linear relationship has also been identified for other products, namely potatoes (Ozuna et al., 2011a) and carrots (García-Perez et al., 2008). Hence, the highest applied ultrasonic power produces the most violent expansions and contractions of the soft solid material and the highest level of internal mechanical stress, thus maximizing the reduction of internal mass transfer resistance. Moreover, microstirring and the generation of microcurrents at the interfaces are more intense and so produce a greater decrease in external resistance.

8.4.1.3 Influence of the Characteristics of the Medium on Ultrasonic Intensity The noncontact applications of ultrasound involve the transmission of acoustic waves across a medium, from the emitter to the product being treated. Due to the fact that ultrasound is composed of mechanical waves, this medium could affect the acoustic field. Media properties such as viscosity, density or surface tension could improve or make it difficult to transmit vibration and, therefore, to extend the ultrasound effects. Figure 8.6 shows the acoustic pressure produced by a sound probe system in saturated brine and in a 30 Brix sucrose solution (Carcel, 2003). As can be observed, when operating under identical conditions, the same ultrasonic system produces a differing acoustic pressure in brine than in a sucrose solution. This means that, even when the same level of electric power is applied to the transducer, the effectively applied ultrasonic energy is different in brine than in sugar solution and (as shown in the previous section) the magnitude of ultrasonic effects will also be different. This fact confirms the suggestions made in Section 8.4.1.1, namely that taking into account only the electrical power supplied/ consumed by transducers constitutes an incomplete approach to the ultrasonic field applied. However, it is not only the physical properties of the medium that have a significant influence on the ultrasonic field applied, but also the level of agitation. Ultrasonic waves are orderly vibrations that are transmitted by interaction between molecules. In the case of agitated fluids, these molecules are moving randomly due to the action of the global fluid movement. This fact hinders the tidy interaction between molecules that is needed to transmit the ultrasonic wave, and transmis- sion of the acoustic wave is thus disrupted. In this sense, Carcel (2003) carried out 8.4 Influence of the Main Process Variables on Drying Intensification by Ultrasound 259

1.1

1

0.9

0.8

0.7

0.6 Acoustic pressure (bar)

0.5 Brine Sucrose solution 0.4 0 20 40 60 80 100 120

Percentage of electric power applied to tranducer

Fig. 8.6 Acoustic pressure measurements probe system (probe diameter 13 mm at carried out with a hydrophone in 800 ml of 15 mm from the emitter surface) supplying saturated brine and a sucrose solution different percentages of the total electrical (30 Brix).Ultrasoundwasappliedwitha power of the equipment (100 W).

ultrasonic field measurements in an ultrasonic bath filled with distilled water. As can be seen in Fig. 8.7, from the measurements carried out without agitation of the bath water it was possible both to identify the stationary field, produced inside the bath by the reflection of ultrasonic waves in the water–air interface, and to locate maximum and minimum acoustic pressure zones. When the water was strongly agitated, the acoustic pressure measured was significantly reduced (to less than one-half), and the differences between maximum and minimum acoustic pressure zones were diminished. It was apparent that the turbulence produced by the water agitation had disturbed the ultrasonic field, in turn reducing the effective ultrasonic energy applied and, therefore, its ability to affect mass transfer processes. The reduction in the ultrasonic field produced by the presence of turbulences occurs not only in liquid but also in gas media. In this sense, García-Perez et al. (2006a) observed that, during the ultrasonic (30 kW m 3) air drying (40 C) of carrot cubes at air velocities in the range 0.6 to 10 m s 1, the effects of ultrasound on drying kinetics depended on the air velocity. At low air velocities (0.6 m s 1), ultrasound application increased the drying rate, but at higher air velocities (>2ms 1) the effect of ultrasound on drying kinetics was diminished and even became negligible. García-Perez et al. (2007) used a diffusion model to quantify the influence of air velocity on ultrasound-mediated effects on the drying rates of carrot, persimmon and lemon peel. In the model used, the external resistance was considered negligible, and therefore the identified effective diffusivity was a lumped parameter that included all effects on both the external and internal 260 8 Drying Assisted by Power Ultrasound

20 Without agitation With agitation Water - air interface

15

10

5 Distance from bottom (cm)

0 0 1 2 3 Acoustic pressure (bar)

Fig. 8.7 Variation of acoustic pressure inside an ultrasonic bath (Fungsonics mod. 28 L, 20 kHz; Fungilab S.A., Barcelona, Spain) containing distilled water (28 l). Measurements were carried out with and without water agitation at different distances from the bottom of the bath.

resistances (Mulet, 1994). The conventionally air-dried samples showed that the identified effective moisture diffusivities increased with the increasing air flow rate. However, this influence was only observed up to a threshold value close to 5ms 1; below this level the process was mainly governed by external resistance. The decrease in external resistance, due to higher air velocities, involved an increase in the effective diffusivity. Above the threshold, water removal was controlled by internal resistance and, for this reason, the effective moisture diffusivity was independent of the air velocity and considered to be a material characteristic. The application of ultrasound increased the identified effective moisture diffusivities, but only at low air velocities. As the improvement observed at high air velocities was negligible, it appears that the influence of ultrasound on drying kinetics disappears when high air velocities are applied. Carcel et al. (2007a) modeled the drying kinetics of persimmon using two diffusion models taking into account (or not) the influence of external resistance to drying. At low air velocities the effective diffusivity and mass transfer coefficient identified from ultrasonically assisted drying were significantly higher (confidence level >95%) than in conventional drying. However, these differences were decreased for both diffusivity and mass transfer coefficient when the air velocity was increased, and became negligible when it was above 4 m s 1. In order to identify the influence of air velocity on the ultrasonic field and, therefore, on the ultrasound effects, the acoustic field inside the drying chamber was measured using a microphone (Riera et al., 2011). When measurements were taken without an air flow, an average sound pressure level of 154.3 dB was obtained for an electrical power applied to the transducer of 75 W (30 kW m 3). However, 8.4 Influence of the Main Process Variables on Drying Intensification by Ultrasound 261

Fig. 8.8 Influence of air velocity on the sound pressure level measured inside a cylindrical transducer. when the air velocity was increased the sound pressure level decreased, and remained almost constant at above 8 m s 1 (Fig. 8.8). This confirmed the fact that turbulences generated by high air velocities disturb the ultrasonic field. The reduction in the actual ultrasonic energy applied to the samples was at least partly responsible for the observed decrease in ultrasound effects and, as a consequence, the application of ultrasound is seen to be more efficient at low air velocities. Currently, certain products must be dried using a low air velocity in order to avoid the appearance of various technical problems, such as casehardening. Moreover, most industrial dryers operate at low air velocities of around 2 m s 1; in these cases, the application of ultrasound could represent a very promising technology, by means of which the kinetics can be intensified while the product quality would remain unaffected. Another important aspect to be considered is the relationship between the applied ultrasonic energy and the mass load. Clearly, when the same level of acoustic energy is applied to dry a large amount of product, the acoustic energy available per unit mass is lower than when a small amount of product is dried, and the capacity of ultrasound to induce effects is also lower. Nevertheless, because the observed effects have thresholds, there is no linear relationship. For this reason, it is important to identify the optimal ratio between the level of applied ultrasonic energy and the mass load of the product treated in order to achieve the most efficient ultrasound application – that is, the greatest effect and, therefore, the lowest energy consumption. Carcel et al. (2011a) studied this issue during the drying of carrot cubes (40 C and 1 m s 1)bothwith(75W, 262 8 Drying Assisted by Power Ultrasound

Fig. 8.9 Identified mass transfer coefficients for carrot drying with (75 W, 21.7 kHz) and without ultrasound applications at different mass load densities. From Carcel et al. (2011a).

21.7 kHz) and without power ultrasound application at several mass load densities (12 to 120 kg m 3). The data were modeled using a diffusion model that took the external resistance into account. In experiments performed without ultrasound application, the increase in mass load density did not affect theeffectivemoisturediffusivity,butdidleadtoasignificant reduction in the mass transfer coefficient. The increase in mass load in the drying chamber created preferential pathways of air flow, which increased the external mass transfer resistance. The application of ultrasound reduced the drying time for all of the mass load densities tested, although increases in the mass load density (over the range studied) did not have any significant effect on the identified effective diffusivities; however, the mass transfer coefficient was significantly reduced (Fig. 8.9). At the highest mass load densities tested (108 and 120 kg m 3), the differences between the mass transfer coefficients of ultra- sonically assisted and conventional drying disappeared. Apparently, the high mass load density prevents the ultrasound from reaching the entire sample surface at an appropriate intensity. Thus, the ultrasound effects on the solid–air interface did not evenly cover the whole solid surface, and consequently the global average effect was decreased. To summarize, it can be stated that the main aspect for consideration when applying ultrasound to intensify drying is not the ultrasonic energy generated, but rather the ultrasonic energy received by the samples. Hence, the higher the energy received the greater the influence of ultrasound on mass transport kinetics. In 8.4 Influence of the Main Process Variables on Drying Intensification by Ultrasound 263 order to increase the amount of received energy, ultrasound application systems must achieve a good coupling between emitters and the treated medium. The attenuation of ultrasonic energy produced by the medium must be minimized, for instance by reducing the turbulence. Finally, an optimal ratio of energy received to mass load treated will allow the greatest ultrasound effects to be achieved with a minimum of energy consumption.

8.4.2 Drying Air Temperature

As in most chemical and biochemical reactions, the temperature can affect the drying kinetics by affecting the moisture diffusivity (Simal et al., 2005). This tendency is usually described by using an Arrhenius-type equation: ! ~ Ea De ¼ D0 exp ; ð8:5Þ RT~ fi ~ where De is the effective diffusion coef cient, D0 is the Arrhenius prefactor, Ea is the molar activation energy for moisture diffusion, R~ is the universal gas constant, and T is the temperature. In overall terms, an increase in temperature accelerates the reactions due to a greater amount of energy being available in the medium, which provides the activation energy needed to “extract” a molecule of water from the product matrix. The application of ultrasound during drying implies the introduction of an additional energy source that could also overcome the barrier of activation energy and make it easy to extract water. Thus, it is expected that ultrasound will have a greater influence on drying when less energy is available in the systems; that is, when drying takes place at low drying air temperature. In contrast, when the drying air provides enough energy – that is, when hot drying air is used – the application of ultrasound has less of an effect, or even no effect at all. Gallego-Juarez (1998), when studying the airborne ultrasound application (155 dB) with a vibrating plate to dry carrots at different temperatures (60, 90 and 115 C), noted that the effect of the ultrasonic radiation was significant at low air temperatures (60 C) but was diminished as the temperature increased. At the highest temperature tested (115 C), the ultrasonic effect was negligible and there were no appreciable differences in the experimental drying curves. In a similar study, García-Perez et al. (2006b) dried carrot cubes (1 m s 1) in a cylindrical vibrating chamber (García-Perez et al., 2006a) at five different temperatures (30, 40, 50, 60, and 70 C), and then used a diffusion model that takes the external resistance to water transport into account to evaluate how temperature can influence the extent of the ultrasound-mediated effects on drying kinetics. The results showed that the effective diffusivity and mass transfer coefficient were both increased as the air temperature rose; this meant that both the internal and external mass transfer resistances had been reduced due to the increase in drying temperature. The influence of power ultrasound (30.8 kW m 3 and 21.7 kHz) on 264 8 Drying Assisted by Power Ultrasound

Fig. 8.10 Influence of drying temperature on the identified effective diffusivity for carrot drying carried out with and without ultrasound application. From García-Perez et al. (2006b).

both kinetic parameters was significant at low temperatures but negligible at the highest temperature tested (70 C). As can be observed in Fig. 8.10, in the case of conventional drying, the dependence of effective diffusivity fitted an Arrhenius-type equation adequately, the activation energy being a characteristic of the product. For ultrasonically assisted dried samples an Arrhenius behavior was observed at low temperatures; however, when the temperature reached 70 C the identified diffusivity was quite similar for both methods of drying, with and without ultrasound application. These facts confirmed that, at low temperatures, ultrasound would provide mechanical energy that complements the thermal energy provided by air temperature and, jointly, these would contribute to the water transport. At high air temperatures, the thermal energy is higher and masks the effect linked to the mechanical energy provided by ultrasound. Therefore, the application of ultra- sound as a means of intensifying air drying processes is more efficient when a low temperature is used. Both, the drying time and drying air temperature determine energy consump- tion, and the application of ultrasound can affect both aspects so as to contribute to a reduction in the energy needs of drying. García-Perez et al. (2012a) used a digital power meter to measure the total energy consumption during ultrasonically assisted drying experiments on orange peel (40 C; 1 m s 1; 0, 18.5, 36.7 kW m 3). The main elements considered for the quantification of energy consumption were the heating elements, the ventilation system (fan), and the ultrasound generator. By applying ultrasound, the drying time was reduced by approximately 30% when the experiments were carried out using an ultrasonic power density of 18.5 kW m 3, 8.4 Influence of the Main Process Variables on Drying Intensification by Ultrasound 265

and by 45% for experiments carried out at 36.7 kW m 3. This shortening of drying times led to a fall in the total measured energy consumption of 12% at 18.5 kW m 3, and of 20% at 36.7 kW m 3 as compared to the conventional drying process. It should be noted that, although these measurements were obtained for a specific dryer, they illustrate the energy-saving possibilities of applying power ultrasound in order to intensify a drying process. As ultrasound application is more effective at low temperatures, this could also represent an interesting approach to intensifying other drying processes that occur at temperatures below or close to the freezing point, such as vacuum freeze-drying or atmospheric freeze-drying. For these drying methods, evaporation or sublima- tion occurs depending on the drying temperature and solvent involved, but long drying times are required in every case. In this context, attempts have been made to apply ultrasound to vacuum freeze- drying processes (Moy and DiMarco, 1972), but with poor results. As ultrasonic waves are mechanical and require a material medium for their propagation, their application under vacuum conditions is difficult. Indeed, in a vacuum only contact ultrasound should be considered. When studying atmospheric freeze-drying, Moy and DiMarco (1970) used a steam-jet whistle operating at frequencies between 10.8 and 12.2 kHz, which is within the human hearing range. Experiments conducted with distilled water, coffee and tea extracts at temperatures of 15 and 26 C revealed an average drying rate increase of between 10% and 100%. However, the intense noise generated by this system would likely be a serious hindrance to its industrial application. Further efforts have been made in this field recently, however, with Bantle and Eikevik (2011) using a commercial transducer (20 kHz; DN 20/2000, Sonotronic) to test the influence of process variables such as temperature and applied power. In these studies, it was shown that ultrasonic application would lead to a maximum reduction of about 10% in the drying time for green peas at 3 C, but that such improvement was almost negligible if the natural variability of raw matter was taken into consideration. A new system by means of which ultrasound can be applied in order to intensify mass transport in low-temperature drying processes has recently been developed (Carcel et al., 2011b). The application of power ultrasound during atmospheric freeze-drying (14 C, 2 m s 1, 0 and 19.5 kW m 3, 21.9 kHz) greatly speeded up the drying kinetics of carrot, apple and eggplant (García-Perez et al., 2012b), and involved an average reduction in drying time of between 65% and 70%. An analysis of experimental data with a diffusion model that takes external resistance into account showed that ultrasound had influenced both the internal and external mass transfer resistances. Thus, the application of power ultrasound had increased the effective diffusivity (Fig. 8.11) in the range of 407% for eggplant and 428% for apple. With regards to the mass transfer coefficient, the increase was 96% for apple, 152% for carrot, and 170% for eggplant. For all of these products such increases in kinetic parameters were significant at a confidence level of 99%. Therefore, the intensifying effect of ultrasound application on effective diffusivity is greater at low temperatures than at high temperatures, above 0 C. In the case of mass transfer coefficients, the increase produced by ultrasound was similar to that 266 8 Drying Assisted by Power Ultrasound

5 30 10 Carrot Eggplant Apple 25 4 8 /s) /s) /s)

2 2 20 2 m 3 m m 6

-11 -11 -11 15 -11 4 (10 2 (10 (10

e e 10 e D D D 1 5 2

0 0 0 AIR AIR+US AIR AIR+US AIR AIR+US

1 Fig. 8.11 Identified effective diffusivities of atmospheric freeze drying (14 C, 2 m s ) of carrot, eggplant and apple, with (AIR þ US) and without (AIR) ultrasound application (19.5 kW m3, 21.9 kHz). The bars indicate intervals with a confidence level of 99%.

reported for conventional hot-air drying (Gallego-Juarez et al., 2007; Puig et al., 2012). It is important to note here that ultrasound effects are more pronounced in regions where the resistance to mass transfer is higher. During freeze-drying, the frozen inner core of the product is shrinking, while a sublimation front is moving back; this generates a porous structure zone between the surface of the material and the front where the internal resistance to mass transfer is located. The effects produced by the application of ultrasound in this zone lead to a significant reduction of this resistance, as observed in the marked increase in identified effective diffusivity. The application of ultrasound at low temperatures is useful not only as a means of intensifying the drying process by sublimation, but also for eliminating other solvents by evaporation. In this context, García-Perez et al. (2012b) monitored the influence of ultrasound in the removal of ethanol from impregnated apple samples at 14 C (previously freeze-dried). The results obtained showed that the time needed to remove all the ethanol fell by 55%, from 150 min for air-drying to 67.5 min for the ultrasonically assisted process. These results demonstrate the potential of this technology for application in the chemical, cosmetic and pharmaceutical industries for the low-temperature removal of organic solvents in order to preserve quality aspects at lower operating and fixed costs than in freeze-drying.

8.4.3 Ultrasound–Sample Interaction

An important aspect to take into consideration when dealing with the application of high-intensity ultrasound as a means of intensifying a drying process is the type of material to be dried and how it might interact with ultrasonic waves. The interaction of ultrasonic waves with different materials could affect the ultrasound efficiency and therefore, a different influence on mass transport might be exerted. Mulet et al. (2003) compared three solid–liquid dehydration treatments (the osmotic dehydration of apple, the brining of cheese, and the brining of pork meat) 8.4 Influence of the Main Process Variables on Drying Intensification by Ultrasound 267 carried out with ultrasound application under similar conditions using a commercial ultrasonic bath. The intensification of mass transport produced by ultrasound was higher in apple than in cheese curds, but no significant ultrasonic effects were noted for meat brining. These different behaviors may relate to the different structures of the products being examined. For example, in apple the structure allows a good transmission of ultrasonic waves inside the product, while the large number of air-filled pores present allow air to flow more easily from the pores to the surrounding solution, and also for the solution to flow inside the pores due to successive compressions and expansions being provoked by ultrasound. In contrast, the “rubber” structure of pressed cheese curds will absorb the ultrasound energy, limiting the transmission of internal vibration; the liquid-filled pores would also complicate mass transfer with the solution, so that the influence of ultrasound would be less than in apples. Finally, at the power level tested, the fibrous structure of pork loin and the absence of pores within the tissues appeared to prevent some of the ultrasonic effects and their influence on mass transfer. When the influence of the intensity of applied ultrasound on ultrasound effects was addressed (see Section 8.4.1), the existence was proposed of an ultrasonic intensity threshold above which the influence of ultrasound on mass transfer would become significant. However, the value of this threshold would depend on the material involved. Subsequently, García-Perez et al. (2009) showed that the influence of power ultrasound on the effective moisture diffusivity in drying carrot showed a threshold of about 15 kW m 3. Nevertheless, due mainly to a wide variability in the raw materials, the difference between the means of ultrasonic and non-ultrasonic treatments for the experiments carried out was not significant (level of confidence >95%) until the acoustic power density in the drying chamber exceeded a value of 20.5 kW m 3 (Fig. 8.12a). This means that, in order to draw effective conclusions from these investigations, a sufficiently large number of experiments must be carried out. Under the same conditions, the effects of ultrasound on the moisture effective diffusivity of lemon peel failed to demonstrate a threshold value, though the mean differences became significant at 12 kW m 3 (Fig. 8.12b). Under similar drying conditions, Ozuna et al. (2010) showed the mean identified potato moisture diffusivity to be statistically different from the not ultrasonically assisted drying (level of confidence >95%) when the acoustic power density applied was higher than 25 kW m 3, while García-Perez et al. (2011) reported a value of 12 kW m 3 for eggplant. The influence of the applied ultrasonic power density on mass transfer coefficient was similar to that observed for effective diffusivity. This means that a minimum level of ultrasonic density must be applied for significant effects to occur, and that this threshold will differ for products with different structures (García-Perez et al., 2008, 2011). The realization that the minimum level of acoustic energy needed in order to observe significant ultrasonic effects on drying depends on the material implies that, under adverse conditions for the transmission of ultrasound (e.g., high air velocities), the material could be prone, or not, to drying intensification by ultrasound. In this sense, based on the modeling of experimental kinetics for carrot, persimmon and lemon peel drying at different air velocities, García-Perez 268 8 Drying Assisted by Power Ultrasound

(a) (b) 1.8 13 Carrot Lemon peel 11 1.6 9 /s) /s) 2 2

m 1.4 m 7 -10 -10 (10 (10 5 De De 1.2 3

1 1 0 80604020 100 0 10080604020 Ultrasonic power density (kW m-3) Ultrasonic power density (kW m-3)

Fig. 8.12 Influence of applied ultrasonic power density on the identified effective diffusivity for the drying of (a) carrot and (b) lemon peel. The bars indicate intervals with a confidence level of 95% (García-Perez et al., 2009).

et al. (2007) noted that the effects of ultrasound on the identified effective diffusivity were significant for all three products at air velocities under 6 m s 1. Above this air velocity, the intensification of ultrasound was negligible for carrot and persimmon, but in the case of lemon peel it was significant up to the highest air velocity tested, of 14 m s 1. This means that, in the case of carrot and persimmon, the disturbing effect that a high air velocity exerts over the acoustic field reduced the energy present in the medium to a level that would not affect the mass transfer process. However, in the case of lemon peel, despite a reduction in acoustic energy, the acoustic level was still able to intensify the drying, confirming that the lemon peel is more sensitive to ultrasonic effects. The level of drying intensification produced by ultrasound also depends on the material treated, this being higher in products with a lower acoustic power density threshold. Table 8.1 shows the identified effective diffusivity for different products dried under similar conditions, with and without ultrasound application. The increase in effective diffusivity varies depending on the product being considered; thus, the application of an acoustic power density of 37 kW m 3 will produce a diffusivity increase of about 40% in drying carrot, and about 200% in the case of eggplant. The same behavior was observed in the mass transfer coefficient (Tab. 8.1); for example, the mass transfer coefficient for carrot was 33% higher when ultrasound was applied, whereas for eggplant the increase reached 230%. These results indicate that those products that present a more compact internal structure (e.g., carrot, persimmon, potato) are less sensitive to the effects of ultrasound on the drying rate. In contrast, the more porous products (e.g., eggplant, 8.4 Influence of the Main Process Variables on Drying Intensification by Ultrasound 269

Tab. 8.1 Identified effective diffusivity (De) and mass transfer coefficient (k) of different products dried (40 C and 1 m s1) with ultrasound (US; 37 kW m3, 21 kHz) and without ultrasound (WUS) application. Incr. (%) is the increase for ultrasonic-assisted (US) compared to conventional (WUS) drying values.

10 2 1 4 2 1 Product De10 (m s ) Incr. k10 (kg H2Om s ) Incr. Reference (%) (%) WUS US WUS US

Carrot 2.0 0.1 2.8 0.3 39 4.6 0.1 6.6 0.5 33.3 García- Perez et al. (2007) Potato 4.6 0.1 7.5 0.3 64 2.0 0.4 3.2 0.4 58.1 Ozuna et al., (2011a) Orange 40.4 4.4 61.3 5.3 52 11.7 0.5 24.3 0.2 107.7 García- peel Perez et al. (2012a) Eggplant 8.9 1.3 27.9 3.0 212 1.9 0.4 6.2 0.9 229.4 García- Perez et al. (2011)

lemon peel) are more prone to mass transfer intensification by ultrasound. Thus, the porosity could be a key parameter with which to define the extent of the effects of ultrasound on mass transport (Carcel et al., 2011a). If the product is highly porous, the transmission of vibrations through the solid matrix would be improved such that the ultrasound could affect the inner part of the samples. Moreover, this structure would facilitate the “sponge effect” because it exhibits a low structural resistance to the mechanical stress. With regards to the influence on the mass transfer coefficient, an external porous layer is more prone to ultrasound effects; hence, microstirring at interfaces, pressure variations, oscillating velocities or the formation of microcurrents can occur inside the external pores, thereby increasing the effective surface area for mass transfer. Ozuna et al. (2011b) were the first to attempt to relate the internal structure of the product with the extent of ultrasound influence on the drying rate. These authors studied the evolution of the effective moisture diffusivity versus the ultrasonic power applied during drying for five very different products, namely carrot, potato, lemon peel, orange peel, and eggplant (Fig. 8.13). As noted in Section 8.4.1.2, above an ultrasonic intensity threshold there is a linear relationship between the ultrasonic power level and the effective moisture diffusivity. For carrots and potatoes, due to the figure scale, this threshold is not highlighted, as in Fig. 8.12. The slope of the linear relationship may be used to estimate the efficacy of the ultrasonic application in that, the steeper the slope the more intense the mechanical effects in the material. These slopes were correlated with instrumental textural measurements of the studied products, thereby identifying a linear relationship between slope and hardness (Fig. 8.14): the higher the hardness of the tested products, the lower the slope. 270 8 Drying Assisted by Power Ultrasound

Fig. 8.13 Linear relationship between effective diffusivity and applied ultrasonic power density for the drying of different products.

Fig. 8.14 Slope of the linear relationship between effective diffusivity and applied ultrasonic power for the drying of different products, versus instrumentally measured hardness. 8.4 Influence of the Main Process Variables on Drying Intensification by Ultrasound 271

Consequently, the mechanical compressions and expansions (“sponge effect”) produced by ultrasound in these materials are more intense in soft products, and this leads to a more effective waterremoval.Inthecaseofthe“chewiness” parameter, a similar behavior is observed. However, other determined textural parameters (cohesiveness, springiness, adhesiveness and resilience) failed to correlate well with the effect of ultrasound. A different situation occurs when ultrasound is applied to frozen samples, for example to intensify the atmospheric freeze-drying process. In this case, the structure and microstructure of the products is changed by freezing. During drying, the loss of moisture by sublimation generates two distinct layers in the solid, an inner frozen core and a highly porous external layer. In this way, regardless of the raw product porosity, the applied drying technique generates a very porous structure and, as a consequence, the ultrasound effects are not located in the ice core but rather in the dry outer layer in which vapor diffusion occurs. Thus, the alternating compressionandexpansioncyclesproduced by ultrasonic waves (sponge effect) will make the diffusion of vapor through this porous solid matrix much easier. The effect of ultrasound on the effective diffusivity is expected to be similar for all products, regardless of their original porous structure. Consequently, when García-Perez et al. (2012b) studied the application of ultrasound during the atmospheric freeze-drying of apple, carrot and eggplant (three products with very different native structures and porosities), they failed to identify any major differences in the increases of effective diffusivities produced by ultrasound. García-Perez et al. (2012b) also showed that the effect of ultrasound on the effective moisture diffusivity was greater (Tab. 8.2) than for air-drying at moderate temperatures (Tab. 8.1). The data in Tab. 8.1 show that, in the case of drying at 40 C, the diffusivity increase produced by ultrasound ranged from 39% for carrot to 212% for eggplant. However, in the case of atmospheric freeze-drying at 10 C the increase was within the range 407–425% for all products. It should also be considered that the acoustic power density applied in the latter case (19.5 kW m 3)waslowerthanthatappliedathigher temperatures (37 kW m 3). The high-porosity matrix generated during freeze- drying in the external layer of the products causes the ultrasound effects to be

Tab. 8.2 Identified effective diffusivity (De) and mass transfer coefficient (k) of different products during atmospheric freeze drying (14 Cand2ms1)withultrasound(US; 19.5 kW m3, 21.9 kHz) and without ultrasound (WUS) application. Incr. (%) is the increase for ultrasonic-assisted (US) compared to conventional (WUS) drying values. Data from García-Perez et al.(2012b).

11 2 1 5 2 1 Product De 10 (m s ) Incr. (%) k 10 (kg H2Om s ) Incr. (%)

WUS US WUS US

Carrot 0.8 0.1 4.2 0.4 425 3.3 1.5 8.3 2.3 152 Apple 1.4 0.7 7.4 2.1 428 4.8 0.2 9.4 0.9 96 Eggplant 4.4 1.7 22.3 4.7 407 23.7 4.3 64.1 10.4 170 272 8 Drying Assisted by Power Ultrasound

more intense than in conventionally hot air-dried products. The high efficiency of ultrasound application to increase diffusion is also linked to the greater acoustic energy absorption in high-porosity products. On the other hand, due to the generation of a porous external layer, the effects of ultrasound on external mass transfer resistance are similar to those reported for the conventional hot-air drying of highly porous materials. As noted above, the effects produced by ultrasound at interfaces may lead to an increase in the effective surface area when they occur inside the external pores of the material.

8.5 Conclusions

The application of ultrasound represents a promising means of intensifying drying systems. The nature of its effects, which mainly are mechanical, could lead to the use of milder drying conditions, thereby preserving product quality. In this respect, the effects of ultrasound are more intense when the drying conditions are more adverse to moisture transport, namely at low temperatures or low air velocities. Under such conditions, the application of ultrasound will increase the drying rate, thereby shortening the processing time. These facts – as well as the relatively small amount of energy needed to generate ultrasound effects – may help to save energy and reduce processing costs. However, in order to achieve a significant influence of ultrasound, it is necessary to generate and transmit the ultrasonic energy to the drying samples in an efficient manner. In addition, as intensification is linked to the acoustic energy that reaches the samples, it is especially important to consider all of the aspects that might reduce the acoustic field, such as the presence of turbulences, the absorption of energy, wave reflections, and mass load distribution. Finally, the internal structure and characteristics of the material to be dried might affect the extent of process intensification by ultrasound. Consequently, porosity would appear to be one of the main factors to be taken into account, as ultrasound is more effective in the most porous products. Clearly, porous structures generated in food products during atmospheric freeze-drying will provide interesting advantages for the successful application of ultrasound as a means of intensifying this process.

Acknowledgments

The authors of this chapter wish to thank the Spanish Ministry of Economy and Competitiveness for their financial support under projects DPI2012-37466-C03-01, DPI2012-37466-C03-02 and DPI2012-37466-C03-03. The authors also acknowledge the contribution of eng. Ramon Pena~ in developing the different drying systems of the ASPA group. References 273

Additional Notation Used in Chapter 8 d distance from emitter surface m f wave frequency s 1 I acoustic intensity kg m4 s 3 (W m 2) k mass transfer coefficient kg m 2 s 1 2 PA acoustic pressure N m Z impedance kg m 2 s 1

Greek Letters

a attenuation coefficient m 1 l wavelength m

Subscripts a activation e effective 0 emitted, reference

Abbreviations

AFD atmospheric freeze-drying d.m. dry matter VFD vacuum freeze-drying

References

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9 Microwave-Assisted Drying of Foods – Equipment, Process and Product Quality Yingqiang Wang, Min Zhang, and Arun S. Mujumdar

9.1 Introduction

Dehydration is one of the main methods used to preserve foods, keeping it in a stable and safe condition as dehydration reduces water activity and extends the shelf life of fresh fruits, vegetables and aquatic products by preventing microbial activity (Wang et al., 2011b; Mujumdar and Law, 2010; Zhang et al., 2006). Furthermore, there are some other objectives of dehydration, for example, quality enhancement, improving ease of handling, further processing, sanitation, and the development of new products (Duan et al., 2010a). Drying is the most common, but also the most energy-consuming, food preservation process (Mujumdar, 2006; Ratti, 2001). Many conventional thermal methods, such as hot-air drying, vacuum-drying and freeze-drying result in low drying rates in the falling rate period of drying and often also in the surface drying region, due to the heat sensitivity of most foodstuffs (Clary et al., 2005; Zhang et al., 2003; Zhang et al., 2005). Furthermore, extended drying times at relatively high temperatures during the falling rate period lead to the undesirable thermal degradation of the product (Mousa and Farid, 2002). In addition, convective drying seriously damages the sensory characteristics and nutritional properties of foods. Occasionally, severe shrinkage caused by convective drying can reduce the rehydration capacity of the dehydrated product. Given the large quantities of materials produced from agriculture, horticulture and other commodities that must be dried each year, and the magnitude of associated energy inputs, there is a continuing need to improve the thermal drying systems used. The main areas for improvement are in energy transfer, drying times to reduce equipment size, and attaining better uniformity of drying rates and product quality (Souraki et al., 2009). Microwave (MW) applications offer potential opportunities for enhancing the drying of food. Microwave heating/drying coupled with other drying techniques such as convective, vacuum- and freeze-drying provides various options for microwave-assisted drying techniques with their respective merits; for example, microwave vacuum-drying (MWVD), microwave freeze-drying (MWFD),

Modern Drying Technology Volume 5: Process Intensification, First Edition. Edited by Evangelos Tsotsas and Arun S. Mujumdar. Ó 2014 Wiley-VCH Verlag GmbH & Co. KGaA. Published 2014 by Wiley-VCH Verlag GmbH & Co. KGaA. 280 9 Microwave-Assisted Drying of Foods – Equipment, Process and Product Quality

Microwave-assisted drying

Microwave Type of drying combined Microwave Combination emission mode with microwaves application time sequence

Microwave-assisted Single-stage Pulsed mode convective drying Microwave combined drying finish drying

Constant mode Microwave-assisted vacuum drying Multistage Complete combined microwave Continuous and drying drying variable mode Microwave-assisted freeze drying

Microwave-assisted spouted bed drying

Fig. 9.1 Classification of microwave-assisted drying techniques.

microwave spouted bed drying (MWSD) and microwave convective drying (MWCD) (Wang et al., 2011b). These microwave-assisted drying processes can also be run under many operating conditions and various energy modes (Arikan et al., 2012). Figure 9.1 shows a number of possibilities of microwave-assisted drying systems that can be exploited. Microwave-assisted drying is based on the volumetric heating of a food material, unlike convective, conduction or radiative heat transfer which is used in conventional drying methods. Microwave heating allows a faster heat transfer in the food material, without the need for large thermal gradients that may damage the product. In the microwave drying of foods, a reduction in drying time of up to 25–90% and a four- to eightfold increase in drying rate, when compared to convective drying, have been reported (Feng et al., 2012). Therefore, microwave drying is a rapid method of moisture removal, in which the energy consumption can be decreased by between 32% and 71% as compared to conventional drying (Duan et al., 2010b; Varith et al., 2007; Durance and Wang, 2002; Sharma and Prasad, 2006a). From the viewpoint of either drying time or drying efficiency, microwave-assisted drying has one major advantage over its conventional drying counterpart in that, for most cases, the quality of the products from microwave- assisted drying under optimal conditions is superior or at least similar to the quality of conventionally dried products (Cui et al., 2003; Lin et al., 1999, 1998). Even microwave-assisted drying such as microwave vacuum-drying can provide dehydrated products that have better characteristics than those dried by freeze- drying due to the development of a desirable, crispy texture in the dried foods, 9.2 Microwave-Assisted Drying of Foods 281 often providing an excellent alternative method to produce no-fried snack products (Wang et al., 2013; Bai-Ngew et al., 2011). Microwave-assisted drying has added advantages that include selective heating, a better and more rapid process control, as well as less floor-space requirements due to the shorter residence time needed for drying (Zhang et al., 2010). However, the drying of food materials may be limited by the danger of burning, due either to temperature runaway (especially when a higher microwave power is used in order to remove moisture more rapidly in the hygroscopic region), or to the development of hot spots, possibly induced by the material’s structure (Holtz et al., 2010). Consequently, both the control of microwave-assisted processes and the selection of a suitable drying strategy for a specific food with the least energy consumption represent crucial problems for the drying of food. It has been reported that large quantities of food materials of different origins and with different technological characteristics, including fruits, vegetables, bean, grain, spice plant, meat, aquatic products, processing waste, restructured products, instant soup, honey, sticky Ganoderma lucidum extract, and pasta, can be dried in microwave-assisted drying equipment. Many nonfood products are also dried in the same equipment, including pharmaceuticals and ceramics.

9.2 Microwave-Assisted Drying of Foods

9.2.1 Basic Principles of Microwave-Assisted Drying

Microwaves are electromagnetic radio waves within the frequency band of 300 MHz to 300 GHz, corresponding to wavelengths of 1 mm to 1 m (Shivhare et al., 2010). Although microwaves cover a wide range of frequencies, their use is restricted to specific frequencies owing to the possibility of their interference with radar or other communication devices. The typical frequencies used in microwaves are 2450 MHz for home-type ovens, and 915 MHz for industrial use (Sahin and Sumnu, 2006). Microwave-assisted drying is based on the volumetric heating by microwaves of dielectric materials such as water, fat and protein, which is a different mechanism from the conduction or convection heating when using an electric heater or flue gas. Conventional heating relies on conduction and convection to transport heat from the heating sources to the product, and this requires a relatively long period of time for heat-sensitive materials and poor thermal conductors such as foodstuffs. In contrast, microwave heating has the potential to deliver heat instantly throughout the product, due to volumetric heat generation. The mechanism of microwave drying involves the exposure of a product to microwaves that penetrate directly into the material, and this results in heating from the inside out. Microwaves are not forms of heat, but rather are forms of electromagnetic energy that manifest as heat through their interaction with the material undergoing exposure (Mujumdar, 2006). 282 9 Microwave-Assisted Drying of Foods – Equipment, Process and Product Quality

Driving force for moisture transport

Temperature Concentration Vapor pressure difference difference difference

Moisture Vapor Hydraulic diffusion diffusion flux

Fig. 9.2 Mechanisms of moisture transport inside unfrozen food materials during microwave heating. Adapted from Cui (2004).

Microwave energy absorption in foods primarily involves two mechanisms: dipolar relaxation; and ionic conduction (Sahin and Sumnu, 2006). In ionic conduction, ions are accelerated by electric fields, causing them to move towards the direction opposite to their own polarity. This movement of ions provokes collisions with the molecules of the material, thereby creating a disordered kinetic energy such that, as a consequence, heat is generated. Polar molecules subjected to microwave radiation at 2450 MHz will rotate 2.45 109 times per second, and the friction between the rapidly rotating molecules causes heat to be generated throughout the material rather than simply be transferred from the surface to the inner part, as is the case for conventional hot-air drying. The water contained in the food serves as the primary dipolar component responsible for dielectric heating (Ratti, 2008), during which different driving forces for moisture transport are developed from the inside to the outside of the heated material. In this way, the water is moved rapidly to the surface by a combination of moisture diffusion, vapor diffusion and hydraulic flow. The mechanisms of moisture transport inside an unfrozen food material during microwave heating are summarized in Fig. 9.2. The moisture, having reached the surface, can be removed by hot air during microwave convective drying, or by a vacuum pump during microwave vacuum-drying or microwave freeze-drying. This restores the driving forces for moisture transport in the heated material. Subsequently, this process is repeated such that the moisture evaporates and leaves the material in continuous fashion, until drying has been completed (Cui, 2004). During this process, a part of the microwave energy is transformed directly to the heat of vaporization. In microwave-assisted freeze drying, the basic principle is similar to that of traditional freeze-drying, but differs from the other microwave-assisted drying processes at the sublimation drying stage. In frozen food materials the moisture is removed by sublimation; consequently, the microwave energy supplies only the heat of vaporization during the sublimation drying stage, but at a faster rate compared to either conventional conduction or radiation (Jangam et al., 2011; Duan et al., 2010b). 9.2 Microwave-Assisted Drying of Foods 283

9.2.2 Energy Absorption by Products During Dielectric Heating

Generally, the average power density P (volumetric absorption of microwave energy, W m 3) produced in a material when exposed to microwave energy is defined by the following equation (Al-Harahsheha et al., 2009):

00 2 P ¼ 2pf e0e E ; ð9:1Þ 12 where f is the frequency (in Hz), e0 is the permittivity of free space (8.86 10 Fm 1), e00 is the effective relative dielectric loss factor, and E is the electric field strength inside the material (in V m 1). Dielectric materials, such as food products, convert electric energy at microwave frequencies into heat. The increase in the temperature of a material due to dielectric heating can be calculated as (Regier and Schubert, 2005):

dT 00 rc ¼ 55:63 10 12fE2e ; ð9:2Þ P dt where cP is the specific heat of the material, r is its density, and dT/dt is the rate of temperature increase (all in SI units). The microwave heating rate depends on dielectric properties of the material.

9.2.3 Dielectric Properties

Dielectric properties are the main parameters that provide information about how materials interact with electromagnetic energy during dielectric heating (Sosa- Morales et al., 2010). Dielectric properties are commonly represented by a complex number, the relative complex permittivity 0 00 e ¼ e je ; ð9:3Þ 0 00 where the real part e is thepffiffiffiffiffiffiffi dielectric constant and the imaginary part e is the dielectric loss factor, j ¼ 1. Here, e0 indicates the material’s ability to store electrical energy, and e00 is a measure of its ability to dissipate the electrical energy in the form of heat (Regier and Schubert, 2005; Wang et al., 2011a). The larger the loss factor, the more easily can the material absorb the incident microwave energy. A material’s dielectric properties are dependent on its moisture content, composi- tion and temperature, as well as on the frequency of the field. Both, the dielectric constant and the loss factor increase with moisture content, but decrease with frequency. Although the dielectric constant increases with temperature, the temperature dependence of the loss factor is unpredictable, so that the loss factor may either increase or decrease with increasing temperature, depending on the range of frequency and the moisture content. The dielectric activity of free water present in the food products is high, whereas bound water exhibits a relatively low dielectric activity (Shivhare et al., 2010). The dielectric constant falls steadily in a 284 9 Microwave-Assisted Drying of Foods – Equipment, Process and Product Quality

Fig. 9.3 Dielectric constant of various food materials at 25 C. Adapted from Sahin and Sumnu (2006).

linear manner as the NaCl concentration increases, while the loss factor increases significantly with the increase of NaCl concentration (Al-Harahsheha et al., 2009). The dielectric constants and loss factors of various food materials are summarized in Figs 9.3 and 9.4, respectively. As can be seen in these figures, the dielectric properties of cooking oils are very low because of their nonpolar characteristics, whereas the dielectric properties of water and high-moisture- containing foods such as fruits, vegetables and meat are high because of dipolar

Fig. 9.4 Dielectric loss factor of various food materials at 25 C. Adapted from Sahin and Sumnu (2006). 9.3 Microwave-Assisted Drying Equipment 285 rotation. The highest loss factor is observed in the case of salt-containing foods such as ham (Sahin and Sumnu, 2006).

9.2.4 Penetration Depth

The penetration depth is an important measure that is often used to assess whether an electromagnetic field at a certain frequency can provide a relatively uniform heating in a given food product (Regier and Schubert, 2005). The penetration depth (dp) is usually defined as the depth into a sample where the microwave power has fallen to 1/e (e ¼ 2.718) or 36.8% of its transmitted value. The penetration depth is a function of e0 and e00, and can be calculated as (Mujumdar, 2006): pffiffiffi l rffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi0 2 dp ¼ qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi ð9:4Þ 2p e0½ 1 þðe00=e0Þ2 1 or pffiffiffiffi l e0 d ¼ 0 : ð9:5Þ p 2pe00 l l ¼ = Here, 0 is the free space microwave wavelength, 0 c0 f , and c0 is the speed of light in free space (c ¼ 3 108 ms 1). 00 Common food products have e < 25, which implies a dp of 0.6–1.0 cm. The penetration of microwaves at 915 and 2450 MHz in foods with high moisture contents at room temperature is typically between 0.3 and 7 cm, depending on their dielectric properties and the frequency of the microwave field (Sosa-Morales et al., 2010). At a constant frequency, the penetration depth increases with the dielectric constant but decreases with the material’s loss factor. Both, the dielectric constant and the loss factor decrease with a reduction in moisture content, but the penetration depth is more sensitive to the variation in loss factor. Hence, the penetration depth increases as the material’s moisture content decreases during processing (e.g., drying) (Shivhare et al., 2010). The penetration depth of food materials is significantly increased as the temperature decreases, especially at temperatures below the freezing point (Bai-Ngew et al., 2011). This is beneficial for the uniformity of microwave heating and, hence, for the improvement of product quality in microwave-assisted freeze-drying (Wang et al., 2011a). The penetration depth falls dramatically as the concentration of sodium chloride in foods increases (Al-Harahsheha et al., 2009).

9.3 Microwave-Assisted Drying Equipment

A microwave-assisted drying system is generally composed of a magnetron as microwave power source, a waveguide, a drying chamber, a material plate, a 286 9 Microwave-Assisted Drying of Foods – Equipment, Process and Product Quality

Fig. 9.5 A simplified schematic of a pilot-scale microwave-assisted convective dryer.

dehumidi fication system, a mechanical drive system, safety protection system, and a sensor and control unit (www.vertpedia.com). Based on the combination of microwaves with other drying techniques, there are mainly four types of microwave-assisted dryer, as discussed below.

9.3.1 Microwave-Assisted Convective Drying Equipment

Figure 9.5 shows a simplified schematic of a pilot-scale microwave-assisted convective dryer described by Pereira et al. (2007) and Ahrne et al. (2007). The microwave dryer consists of a microwave generator and a forced air system. The microwave generator has a maximal microwave power of 1000 W operating at a frequency of 2450 MHz. The cylindrical cavity of the dryer with a vertically oriented axis has a diameter of 80 cm and a height of 10 cm, and accommodates a Teflon tray and a built-in scale with an accuracy of 2 g. The weight loss, product temperature and microwave power were recorded every second by a data acquisition system. The air flow inlet and outlet are located in the middle of the chamber, and the heated air is distributed and directed through the lid to the samples placed inside the cavity. Microwave power, air temperature, velocity and humidity were adjustable parameters. The product temperature was measured by fiber-optic temperature sensors (ReFLexTM, Neoptix, Canada) placed in four samples located at different positions of the sample plate. The size of the cavity was optimized by modeling the microwave field distribution for 1 kg of fruit pieces distributed as a single layer. In this dryer, the load does not rotate. A laboratory-scale microwave convective dryer with maximum output of 625 W at 2450 MHz and cavity (drying chamber) dimensions of 300 mm (width), 240 mm 9.3 Microwave-Assisted Drying Equipment 287

(depth) and 210 mm (height) was reported by Uprit and Mishra (2003). The dryer had the capability of online sample weighing and power (wattage) adjustment, and the load could rotate, but the material temperature was not controlled due to the movement of material. There are three methods in which microwave energy may be combined with hot- air drying, by applying the energy:

at the beginning of dehydration processes; when the drying rate begins to fall; or during the falling rate period(s) or at low moisture content to facilitate finish drying. In some cases, applying microwave drying at the last stage of the dehydration process can also be very efficient for removing bound water from the product (Zhang et al., 2006).

In the microwave-convective air-drying systems, microwave energy spurs the water to reach the surface of the product at a high rate, while the convective flow of air helps in removing the moisture from the drying chamber and also contributes to a more homogeneous drying. Thus, combined microwave and convective drying systems increase not only the drying rate of the product but also the quality of the dry product obtained.

9.3.2 Microwave-Assisted Vacuum Drying Equipment

Figure 9.6 shows a laboratory microwave-vacuum dryer on the basis of hexahedron of the cavity with the dimensions of 360 mm 340 mm 260 mm. The designed nominal power output of the magnetron was 750 W, and the actual measured output was 400 W. The turntable and the motor were connected by a shaft which was dynamically sealed. The structure and dimension of the window for the microwaves to pass was carefully designed in order to avoid any electrical discharge due to lower pressure. The door was sealed by a

5 6 4 7 8 3 9 2 1

Fig. 9.6 Schematic of a laboratory microwave-vacuum dryer. 1, Vacuum pump; 2, Rotating tray; 3, Valve; 4, Observation window; 5, Drying chamber; 6, Infrared thermometer; 7, Vacuum gage; 8, Magnetron power control unit; 9, Magnetron. 288 9 Microwave-Assisted Drying of Foods – Equipment, Process and Product Quality

Fig. 9.7 A pilot microwave-vacuum dryer.

convex–concave face with a microwave leakage of less than 5 mWcm 2.The microwave-vacuum dryer was operated at 25 mbar and the rotation speed of the turntable was 5 rpm (Cui et al., 2003). Figure 9.7 presents, schematically, a pilot vacuum microwave dryer (WZD4S-1, Nanjing, China) with a maximum nominal power of 4.2 kW and an operation frequency of 2450 MHz. The operation power is adjustable between 700 W and 4200 W at 700 W increments. The dimensions of the microwave cavity were 1050 mm, 1080 mm, 800 mm (width, depth and height, respectively). The minimal attainable pressure is 5 kPa. Six material plates in the cavity can move around at 1 rpm during operation. Figure 9.8 shows a schematic diagram of a microwave-vacuum rotary drum dryer reported by Kaensup et al. (2002). In this dryer, a cylindrical vacuum chamber was made from stainless steel having diameter 0.33 m and length 0.4 m. A 2450 MHz magnetron having maximum power output of 800 W was located 110 mm aside of the drying chamber. Microwaves were conducted to the vacuum chamber via a stainless steel waveguide having a rectangular cross- section of 40 mm 80 mm. Polypropylene was chosen to make a perforated rotating drum with guide vane, as this material is almost transparent to microwaves. The drum had a diameter of 0.3 m and length of 0.3 m. A novel vacuum seal made from metal shield attached on Teflon bush enabled the rotary drum to be driven in the chamber by means of an external gear motor. The enclosed chamber was almost microwave-proof, with measured leak radiation at 0.5 mWcm 2.Avacuumpumpwithacapacityof63m3 h 1 at 4 mmHg was used to withdraw vapor from the vacuum chamber. The drum rotation speed was controlled using an electronic inverter. Electrical energy 9.3 Microwave-Assisted Drying Equipment 289

Fig. 9.8 A microwave-vacuum rotary drum dryer. Adopted from Kaensup et al. (2002).

consumed in the drying process was measured by means of a Watt–hour meter. In operation, the guide vanes agitate the product as the drum is rotated in the chamber, which allows for an even exposure of the product to microwave energy and a more uniform mixing. Furthermore, the moisture expelled during the drying process can easily leave the product, resulting in a shortened drying time and no hot spots. The application of an industrial-scale microwave-vacuum dryer (Model MG8KW, Enwave Corporation, Vancouver, Canada) to dry Saskatoon berries was reported by Pranabendu and Venkatesh (2009). The dryer consisted of a cylindrical stainless steel chamber, where the microwave field and vacuum were applied. Inside the chamber, the berries were contained in a basket made from ultrahigh-density polyethylene, sitting on two rollers made from the same material. One of the rollers was connected directly to a variable-speed motor through the chamber wall; the cogs in the wall of the basket mesh matched the cogs in the roller to provide a smooth rotation. The basket rotation speed was set prior to each drying cycle and kept constant at 10 rpm for all of the drying treatments. The vacuum was provided by an external pump, and the vacuum pressure was kept constant at about 100 kPa (Pranabendu and Venkatesh, 2009). The recommended optimum drying condi- tions for the drying of Saskatoon berries are a microwave power of 5.7–6 kW and a drying time of 51.5–55 min, with a fruit load of 9.75 to 10 kg.. 290 9 Microwave-Assisted Drying of Foods – Equipment, Process and Product Quality

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18

19 20 T 21

Fig. 9.9 Laboratory-scale microwave-assisted Heating plates; 11, Freeze-dried sample; freeze-drying equipment. 1, MWFD 13, Pressure senor, for freeze-drying chamber; 2, Fiber-optical temperature chamber; 14, Vacuum breakage valve, for sensor; 3, Vacuum breakage valve, for freeze-drying; 15, Temperature sensor; 16, MWFD; 4, Sample supporting plate; Freeze-dryer vacuum valve; 17, MWFD 5,MWFDsample;6and7,Microwave vacuum valve; 18, Cold trap; 19, Vacuum sources; 8, Pressure senor, for MWFD pump; 20, Draining valve; 21, Refrigeration chamber; 9, FD chamber; 10 and 12, compressor; 22, Control system.

9.3.3 Microwave-Assisted Freeze-Drying Equipment

A typical schematic diagram of microwave-assisted freeze-drying equipment is shown in Fig. 9.9. During drying, the pressure was maintained at 100 Pa by a vacuum pump, and the temperature of the cold trap (40 to 45 C) was sufficiently low to condense the vapor. The microwave frequency was 2450 MHz and the power could be regulated continually from 0 to 2000 W. Amicrowavefield can be used as the heat source to supply the heat of sublimation needed in the freeze-drying process. Volumetric heating of the material in the vacuum environment makes it possible to greatly improve the rate of sublimation. In the equipment, the microwave resonant cavity was designed as an effective multimode resonant cavity, which makes the electric field distribution uniform when the material to be dried is in a stationary state (Duan et al., 2010b; Wang et al., 2011b). It should be noted that subdivision of the drying chamber of Fig. 9.9 to two parts enables microwave-assisted freeze- drying and conventional freeze-drying to be conducted in parallel, under otherwise the same conditions. 9.3 Microwave-Assisted Drying Equipment 291

Fig. 9.10 Schematic diagram of microwave- window; 10, Operation door; 11, Spouted assisted spouted bed. 1, Air compressor; bed; 12, Air flow detector; 13, Control box; 2, Blower; 3, Electric heating; 4, Pressure 14, Power regulator; 15, Optical fiber detector;5,Valve;6,Nozzle;7,Peristaltic measurement; 16, Equipment support. pump;8,Magnetron;9,Observation

9.3.4 Microwave-Assisted Spouted Bed Drying Equipment

Figure 9.10 shows a typical schematic diagram of a microwave-assisted spouted bed. The system consisted of 2450 MHz microwave power sources, a cavity, a hot- air source, a spouted bed, and a control system. Microwave power was provided by four magnetrons, each having 1 kW maximum power capacity. The total microwave power could be regulated between 0 and 4 kW. The temperature of the air for the spouted bed could be controlled between 30 and 150 C through a feed-back loop, and the air velocity was maintained with an adjustable fan. A fiber-optic temperature sensor detected the temperature in the drying cavity. The bottom of the cone was made from a plastic screen to hold the particulate samples and provide a pass for hot air. A metal screen with small holes was placed on the observation window to stop microwave leaking. Figure 9.11 shows the scheme of a laboratory-scale microwave-assisted inert medium fluidized-bed drying experimental apparatus. A domestic microwave oven (Type: LG, MC-2003TR(S)) with a frequency of 2450 MHz was used to set up the dryer. By making a hole under the microwave oven and sealing it, a fluidized bed was placed into the microwave oven. A hole was also made on top of the microwave oven so that air could exit. U-shape metallic tubes were used at the inlet and outlet of the drying air stream to prevent microwaves from coming out of the oven. The other parts of the system were the same as those used in a conventional fluidized- bed dryer. The dryer was a cylindrical Pyrex column of 77 mm diameter and 35 cm 292 9 Microwave-Assisted Drying of Foods – Equipment, Process and Product Quality

Microwave oven

Timer Inert particles Power set switch Drying Temperature sample Rotameter Controller Air distributor

Air Flow rate regulating valve Electrical heater

Fig. 9.11 Scheme of microwave-assisted inert medium fluidized-bed dryer system. Reproduced with permission from Souraki et al. (2009).

height, equipped with a porous plate for air distribution. An overall amount of 400 g of glass beads of diameter 2.7 mm was used as energy carrier. Although the glass beads are transparent against microwaves, they increase the rate of heat and mass transfer, and act similarly to the turntable in conventional domestic microwave ovens. The fluidization of inerts agitates the drying sample (carrot), so that microwave radiation can be received by all of the sample parts. Drying air was supplied from a high-pressure air source, with the pressure being adjusted by a regulator. Air was passed through a rotameter and then heated by a controlled electrical heater. A temperature controller was used to regulate the temperature of drying air within 1 K. Air humidity was determined by measuring the dry and wet bulb temperatures (Souraki et al., 2009). To date, only microwave hot-air dryers used principally in finish drying processes to level-off the moisture content in pasta, crackers or chips, and microwave vacuum dryers custom-made to specific product requirements, are available commercially (Ratti, 2008). Good microwave-assisted drying equipment should have a uniform distribution of the microwave field, provide a movement of the material to be dried, keep the microwave field strength in the drying chamber lower than the dielectric breakdown strength, and possess an economic dehumidification system.

9.4 Microwave-Assisted Drying Process

Moisture and temperature history relate to quality changes during microwave- assisted drying processes and have, therefore, been intensively investigated. Their 9.4 Microwave-Assisted Drying Process 293 observation and modeling is helpful for understanding and controlling the drying process as well as for process optimization.

9.4.1 Moisture Loss

As drying proceeds, the moisture is removed continuously from the food material. Moisture loss is generally described by the drying rate (in kg of removed water per kilogram of dry material and time), and by the total drying time spent to remove a specific amount of water in a dryer. In general, a complete microwave-assisted drying process displays three drying periods: a transient rising rate period; a rapid rate drying period; and a falling rate drying period. At the start of drying, the drying rate rises rapidly due to fast microwave heating, until the process enters into the rapid rate drying period, when most of the free moisture is removed at a constant rate or a high falling rate. The drying process then continues to its end at low falling rates (see Fig. 9.12; cf. also Wojdyo et al., 2009; Figiel, 2010; Arikan et al., 2012). The drying rate in the rapid drying period is governed mainly by the rate of external heat and mass transfer, since a film of free water is always available at the evaporating surface. The duration of the rapid rate drying period is dependent on the level of microwave power used, the type of material being dried, and the other drying conditions. As the water content is decreased markedly towards the final drying, two parallel types of behavior can be identified: (i) a considerable decrease in the microwave power absorbed by the foodstuff (associated with a smaller e0 and e00); and (ii) a higher resistance to water transport as a result of dried cell layers and a reduction in the driving force caused by concentration and pressure gradients. Both effects contribute to further decreasing the drying rate (Pereira et al., 2007). Compared to its conventional counterpart, moisture can be removed at higher rate during microwave-assisted drying, as shown in Fig. 9.12.

Fig. 9.12 A comparison of typical drying curves between microwave-assisted drying and conventional drying. 294 9 Microwave-Assisted Drying of Foods – Equipment, Process and Product Quality

Factors affecting the drying rate include the type of conventional drying combined with microwaves, microwave power, pulsed mode, load, pretreatment, sample thickness and composition, as well as the performance of the dehumidifica- tion system. The drying rate significantly increases and the drying time falls as the microwave power, or the sum of microwave power over time in pulsed mode, is increased. The drying time increases with an increase in the load (Hu et al., 2006; Esturk, 2012; Wojdyo et al., 2009; Figiel, 2010; Arikan et al., 2012). Osmotic treatment reduces the initial moisture content and increases moisture mobility within the sample by diminishing shrinkage, thus significantly reducing the subsequent drying time (Wang et al., 2010a; Al-Harahsheha et al., 2009). The addition of salt and sucrose to restructured potato slices and instant vegetable soup led to a significant increase in the drying rate, due to an enhanced ability of microwaves to be absorbed by these materials during microwave-assisted freeze drying. It was noted that a 17–54% reduction in drying time occurred as the salt content was increased from 1% to 7%, while a reduction in drying time of between 11.1% and 33.3% was found as the sucrose content was increased from 3% to 15% (Wang et al., 2011a; Wang et al., 2010b). With regards to the effect of sample thickness on drying time, it has been concluded by some groups that the drying time when using the microwave method becomes longer for thinner slices (in contrast to the convective method). For example, the use of thinner banana and potato slices, as well as the slicing of garlic cloves, prolonged the duration of microwave-assisted drying (Maskan, 2000; Wang et al., 2004; Figiel, 2009). This may be attributed to the smaller thicknesses of material reducing the absorption of microwave energy if the penetration depth is too large (Figiel, 2009). Others have reported the same trend in results as for conventional convective drying method, notably in the cases of microwave-assisted freeze-drying of instant vegetable soup and microwave-assisted vacuum-drying of mushroom (Giri and Prasad, 2007a; Wang et al., 2009). In addition, during microwave-assisted convective drying, air (both temperature and velocity) plays an important role, not only as a carrier of evaporated moisture but also as it contributes to a more homogeneous and faster drying (Pereira et al., 2007; Feng and Tang, 1998; Funebo and Ohlsson, 1998; Dev et al., 2011). The air during the initial drying stage does not have any significant effect on the drying rate of microwave-assisted spouted bed drying, but it does have an important effect on the final stage of drying (Feng and Tang, 1998). It was reported that the increase in air velocity in the microwave convective drying of banana from 0.3 to 0.8 m s 1 caused a 20–25% reduction in the drying time (Ahrne et al., 2007). During microwave-assisted vacuum drying, a higher vacuum increases the driving force for mass transfer and also facilitates the evaporation and volatilization of water from the materials, thus shortening the drying time. However, the effect of system pressure on drying time was not as significant as that of microwave power (Hu et al., 2006). Typical drying curves of three different microwave methods are shown in Fig. 9.13. There are insignificant differences in the shape of the drying curves, but the drying time necessary to reach the desirable moisture content differs among 9.4 Microwave-Assisted Drying Process 295

Fig. 9.13 Drying curves of three microwave-assisted drying methods.

the various microwave-assisted drying methods (Yan et al., 2010). The drying time of MWFD was 3, 4.3, and 6 times longer than in MWVD, MWSD at 2.0 Wg 1 and MWSD at 3.5 Wg 1, respectively. The low drying rate in MWFD could be explained by the fact that the dielectric constant value of ice is lower than that of liquid water, which results in a lower conversion of microwave energy to thermal energy in MWFD. In addition, the high drying rate in MWSD can be attributed to the constant movement of carrot pieces within the microwave cavity, which allowed the different parts of the carrot pieces to receive a relatively uniform microwave radiation over a period of time. Table 9.1 shows some drying kinetic model equations used to fit microwave drying data.

9.4.2 Temperature Distributions

Temperature was found to be the most significant predictor of drying process during microwave-assisted drying, as it reflects the general drying performance rather well (Duan et al., 2010b). Three different temperature measurement systems are available (Regier and Schubert, 2005):

Infrared temperature measurement Fiber-optical measurement Metal-shielded thermocouples.

The infrared thermometer can record the surface temperature of the materials, while the metal thermocouples and the fiber-optical temperature 296 9 Microwave-Assisted Drying of Foods – Equipment, Process and Product Quality

Tab. 9.1 Some drying kinetic models for microwave-assisted drying of food (symbols are explained in the text and chapter notation; all other symbols represent fitting parameters).

Drying Model Sample method

MWVD MR ¼ expðktnÞ Mushrooms (Giri and Prasad, 2007a) MWCD MR ¼ expðktnÞ Moringa oleifera (Dev et al., 2011) MWCD MR ¼ exp½ðÞaPMW þ bT t Cooked chickpeas (Gowen et al., 2006) MWCD MR ¼ expðÞktn Cooked soybeans (Gowen et al., 2008) MWVD First period: X ¼ A Bt Strawberry fruits (Wojdyo et al., First period: X ¼ C þ ½D expðÞEt 2009) P t MWVD : ¼ abs Carrot (Cui et al., 2004) X 2 X Xi D Mdry hv  0:5741 Pabst 0 < X < 2 : X ¼ 1:3662X Xi MdryDhv ¼ þ b MWVD X a þ ½ðÞ = Garlic (Figiel, 2009) 1hiexp t c d b MWCD MR ¼ exp ðÞt=a Osmotically treated pineapple (Botha et al., 2012)

sensor can measure the interior temperature. An infrared thermal imaging camera (IRI 4010 Multi-Purpose Imager; IRISYS, UK) was used to determine the temperature distributions of individual thin chips and to assess hot spots. It is reported that for a sample thickness less than 8 mm, the core temperature of thesamplewasthesameasitssurfacetemperature,withuniformtemperature distribution within the sample. However, for sample thickness more than 8 mm, a temperature gradient developed along the thickness of the sample. Generally speaking, if foods such as fruits and vegetables are cut into slices with thickness less than 10 mm before drying, then the sample can be considered to be of uniform temperature (Cui et al., 2005). Figure 9.14 presents typical temperature curves of microwave-assisted drying for food. Here, the shape of the temperature curve of microwave-assisted freeze-drying can be seen to differ from that of the three other types of microwave-assisted drying, due to different drying mechanisms. Temperature changes are heavily dependent on the microwave power, microwave emission mode, type of dryer used, and the presence or absence of a temperature control unit.

9.4.2.1 Temperature Variations at Fixed Levels of Microwave Power Three different temperature stages are observed at constant microwave power in most microwave-assisted drying processes (Clary et al., 2005; Feng et al., 1999; Ahrne et al., 2007; Duan et al., 2010b): 9.4 Microwave-Assisted Drying Process 297

Fig. 9.14 Typical temperature curves of at constant microwave power; 2, Adjusted microwave-assisted drying for food, where 1 power with feedback temperature control, or and 2 indicate the temperature changes in variable microwave power, or pulsed mode, the final drying stage. 1, Convective, or microwave-assisted spouted bed drying vacuum or freeze microwave-assisted drying under constant microwave power.

A warming-up period in which the product temperature increases linearly with drying time. A constant temperature period, which corresponds to the aforementioned constant drying rate period in which moisture is removed at the outer surface of the material and the cooling effect of vaporization remains in balance with the microwave heating. A heating-up period, in which the drying rate decreases and the sample temperature increases rapidly.

During the period of relatively constant temperature, the temperature of microwave-assisted freeze-drying is lowest, which should be lower than the co- melting temperature of the material, while the temperature of microwave- assisted vacuum-drying is generally between 26 Cand52C, depending on the operating pressure. Microwave-assisted convective drying can be run at a higher temperature, depending on the microwave power used and other conditions such as air temperature (Pereira et al., 2007). Similar to the duration of the rapid rate drying period, the duration of the constant temperature period depends on the level of microwave power used, the load, the nature of the material, as well 298 9 Microwave-Assisted Drying of Foods – Equipment, Process and Product Quality

as the other drying conditions. When most of the free water has been removed, the temperature rises sharply (as shown in Fig. 9.14); this is known as temperature overrun or thermal runaway, due to both an increase in the specific microwave energy input (the energy delivered per unit mass of wet product) and a decrease in the product’sspecific heat capacity (this is involved in the conductive heat transfer) with progressing drying (Figiel, 2010; Botha et al., 2012). The microwave power level has a decisive effect on the thermal runaway, with a higher microwave power resulting in a more serious thermal runaway. Low microwave power, and air with a low temperature and a high velocity, had a lowering effect on the thermal runaway during microwave-assisted convective drying (Pereira et al., 2007). In the microwave-assisted spouted bed drying of diced apples, no rise in product temperature was observed with a microwave input power of 6.4 Wg 1 (dry basis), but the temperature of the diced apples was 12–15 C above the spouted bed air temperature. The reason for this may be the high surface heat and mass transfer rates, facilitated by the intensive pneumatic agitation, which ensure heat removal at a high rate and enough convective cooling to balance the heat generated internally by microwave energy. The MWSD system provided a unique temperature-leveling effect after a short warming period (Feng et al., 1999). The results of Figiel (2009) showed that the temperature change in garlic cloves and slices subjected to MWVD can be approximated with the exponential function: T ¼ 50:6 þ 25:9 expðÞX=0:167 : ð9:6Þ According to a model developed by Cui et al. (2005), the temperature of sliced carrot that has been dried in a microwave-vacuum dryer can be predicted by the following equations:

Pabst Warming-up period; X ¼ 7:68: T ¼ T i þ Mwetcp ð9:7Þ Constant-temperature period; 2 X < 7:68: T ¼ T ; < < : ¼ 0:0196 0:0797 Heating-up period 0 X 2 T Pabs X T Here, T is the saturation temperature of water at the prevailing pressure, and Pabs is the absorbed microwave power (in W). The equations were derived for mass- specific microwave power densities from 3.5 Wg 1 to 18.7 Wg 1.

9.4.2.2 Temperature Variations at Variable Microwave Power without Controlling Temperature Variable microwave power means that the microwave energy is emitted in a pulsed mode by turning on-off the microwave power during the drying process, or that the microwave emission is declining continually with time. Compared with the constant power mode, the variable microwave power mode is able to stabilize the temperature of the sample at a value that depends on the microwave power used, the duration of the power-on and power-off periods, and the other operating conditions, which indicates that thermal runaway can be inhibited in the variable microwave power mode. 9.4 Microwave-Assisted Drying Process 299

It was reported that intermittent microwave-convective drying conducted at a low drying air temperature and microwave power level with a relatively long power-off time, resulted in a more stable and gentle drying process due to decreasing the temperature (Soysal et al., 2009; Changrue et al., 2008). Botha et al. (2012) found, in microwave-assisted convective drying of osmotically treated pineapple, that when a declining mode with time was used the drying temperature could be effectively controlled due to the maintenance of a relatively constant specific microwave energy input (the energy delivered per unit mass of wet product) during drying. High microwave powers need to be reduced quickly, and faster than the decrease in water content, in order to minimize charring. An inlet air temperature of 70 C was found to be excessive when combined with microwave energy, and resulted in a rapid temperature increase.

9.4.2.3 Temperature Change with Time-Adjusted Power in Feedback Temperature Control The best temperature control is achieved with a phase controller, which can continuously and automatically adjust the variable microwave power, based on predefined power profiles (Li et al., 2010). It is worth emphasizing here that maximum power levels must be selected, based on trial and error, for a designated temperature. If the power was too high, the temperature fluctuations were too large, but if the power was too low, then during certain drying stages the temperature could not reach the preset temperature. Li et al. (2010) investigated the microwave-assisted convective drying of apple slices with time-adjusted power in feedback temperature control. Three desirable temperatures of 75, 65, and 55 C were chosen, and the corresponding maximum power requirements were set at 400, 300, and 240 W, respectively. The relationships of power versus time and of power versus moisture content (d.b.) can be described by the following equations, respectively: "# Xi¼3 2 t bi PtðÞ¼ ai exp ; ð9:8Þ i¼1 ci "# Xi¼3 2 X ei PXðÞ¼ di exp ; ð9:9Þ i¼1 f i where ai to fi are temperature-dependent constants with appropriate units.

9.4.3 Energy Consumption

During microwave-assisted drying (Varith et al., 2007; Holtz et al., 2010), the energy consumption comprises of:

the energy used for producing microwaves; the energy consumed by the dehumidification system, for example, for heating air or maintaining vacuum and refrigeration; and 300 9 Microwave-Assisted Drying of Foods – Equipment, Process and Product Quality

the energy consumed by mechanical drive systems for, for example, the fan and turntable.

The electrical consumption can be recorded using a Watt–hour meter. The energy consumption of microwave-assisted drying can be evaluated (Varith et al., 2007; Soysal et al., 2006) by two different efficiency indices, namely specific energy consumption (SEC): 3600E SECðÞ¼ kJ=kgH2O ð9:10Þ MsðX i X f Þ

where E is the total electrical energy consumption (in kWh), X i and X f refer to the initial and final moisture contents (d.b.), respectively, and Ms is the mass of dry solid (in kg), or energy efficiency: ÀÁDh gðÞ¼% M X X v 100 ð9:11Þ s i f 3600E 1 where Dhv is the evaporation enthalpy of water (2257 kJ kg , at 100 C). The typical variations of specific energy consumption and energy efficiency with moisture content during microwave-assisted drying are illustrated in Fig. 9.15. At the start of the drying process the specific energy consumption is high and the energy efficiency is low, as the microwave energy applied is used to raise the material’s temperature and very little moisture is evaporated. Subsequently, a decrease in specific energy consumption and a sharp increase in energy efficiency are observed, and both of these maintain a relatively constant level during drying over a wide range of moisture contents. The specific energy consumption is approximately the same for all materials at intermediate to high moisture contents. When drying continues into the hygroscopic region (lower than a critical moisture

Fig. 9.15 Variation of specific energy consumption and energy efficiency with moisture content during microwave-assisted drying. 9.4 Microwave-Assisted Drying Process 301 content), where water is more tightly bound, the specific energy consumption increases while the energy efficiency reduces (Holtz et al., 2010). At lower moisture levels, the absence of water causes e00 to fall considerably, and the material is now much less inclined to absorb the offered energy, thereby causing the efficiency of the dryer to drop (Kaensup et al., 2002). The energy consumption of microwave-assisted drying for food is affected by drying time, microwave density, microwave mode, load, the drying method combined with microwaves and the type of combination, as well as other drying conditions such as air temperature, air velocity, pressure, and drum rotating speed. The specific energy consumption is significantly decreased with an increase in microwave density (Sharma and Prasad, 2006a). Pulsed or intermittent microwave applications can result in a more energy-efficient drying than that in continuous microwave mode (Clary et al., 2007; Sunjka et al., 2004; Arikan et al., 2012). In both the pulsed and the continuous cases, the drying efficiency was improved when a lower pressure or temperature was applied (Yongsawatdigul and Gunasekaran, 1996). Sunjka et al. (2004) found the microwave-assisted vacuum drying of cranberries to be more energy-efficient than the microwave-assisted convective drying of the same product. The efficiency of microwave-assisted convective drying was increased in microwave mode with a longer power-off time, but was decreased in the corresponding microwave mode of microwave-assisted vacuum drying. A lower material load tended to decrease the drying efficiency and to increase the specific energy consumption. About a 9.5% increase in drying efficiency and about fi an 18% (0.92 MJ per kg H2O) decrease in speci c energy consumption could be obtained by working with a 128.6 g material load instead of 64.3 g. As the intensity of heat generation is proportional to the content of moisture in a dielectrically dried material, a larger amount of water trapped inside the material would provide a higher drying efficiency and lower specific energy consumption values (Soysal et al., 2006). It was reported that varying the drum rotation speed and the vacuum pressure had no effect on specific energy consumption in the high moisture content range. Rather, the specific energy consumption was insignificantly changed when the moisture content in the product was high for all drum rotating speeds and chamber vacuum levels investigated. However, for a low moisture content, lower drum rotating speeds and a lower chamber pressure resulted in less energy consumption (Kaensup et al., 2002). It was observed during the microwave-assisted convective drying of garlic cloves that, as the air velocity was increased from 1.0 to 2.0 m s 1 at a given air temperature of 70 C, the specific energy consumption also increased from 26.3 to 63.0 MJ kg 1. This was because the increase in air velocity resulted in a cooling of the drying product; that is, its temperature was reduced such that the drying time was increased as a consequence (Sharma and Prasad, 2006a). When Yan et al. (2010) compared the energy consumptions of different microwave-assisted drying processes (Fig. 9.16), MWVD and MWSD were found to have about 75% less energy consumption than MWFD. Moreover, there was no significant difference in energy consumption between MWVD and MWSD. The difference from MWFD was attributed to the fact that the drying time for MWFD 302 9 Microwave-Assisted Drying of Foods – Equipment, Process and Product Quality

Fig. 9.16 Comparison of the energy consumption of different drying methods. Note: Different letters indicate a statistically significant difference.

was long (approximately five times longer than for MWVD and MWSD) and more energy was needed to maintain a high vacuum and a low temperature in the cold trap. Table 9.2 presents the SEC and drying efficiency values for different microwave- assisted drying processes of food materials, as determined by various research groups. The minimal theoretical value of SEC for a dryer is 2.3 MJ per kg of evaporated water, based on the maximal theoretical specific moisture extraction ratio (SMER) of 1.55 kg per kWh and the latent heat of water evaporation at 100 C (Holtz et al., 2009). These results are expected to depend heavily on the scale and design of dryers, as well as on the method used to measure energy consumption (Durance and Wang, 2002).

9.4.4 Dielectric Breakdown

Dielectric breakdown occurs when the electric field intensity in the drying chamber is above a threshold value, and leads to burning of the product when food materials are being dried. Ionization of the gases present in the drying chamber, such as vapor and air, leads to the appearance of purple light, which causes burns on the product surface. The occurrence of this phenomenon leads to major energy losses and excessive heating of the dry zone of the material, seriously damaging the final product. The threshold value of the electric field is normally a function of chamber pressure. The dielectric breakdown easily occurs when the pressure is about 133 Pa – a value commonly used in conventional freeze-drying operations – and when the moisture content of the 9.4 Microwave-Assisted Drying Process 303

Tab. 9.2 Specific energy consumption (SEC) and energy efficiency for microwave-assisted drying processes of food materials.

Drying SEC Energy Material; Varied Reference 1 process (MJ kg H2O) efficiency (%) parameters

MWCD 4.2–5.6 44.2–53.7 Parsley; load Soysal et al., 2006 MWCD 15–18 15.3–12.8 Potato, bread Holtz et al., 2010 MWCD 10–8.3 23–26 Cranberries; Sunjka et al., 2004 Microwave density and mode MWCD 36.4–29.7 7.7–6.3 Longan; Varith et al., 2007 Microwave density, air temperature MWCD 26.3–62.0 8.7–3.7 Garlic cloves; Sharma and Prasad, Air velocity, microwave 2006b density MWVD Beginning 11.5–3.8 Chili; Pressure, Kaensup et al., 2002 period: 60–20; rotational speed of drum Constant rate 11.5 period: 20; Final drying 11.5–0.32 period: 20–700; MWVD 2.7–4.9 86.5–46.9 Cranberries; Yongsawatdigul and Pressure, microwave Gunasekaran, 1996 power, microwave mode MWVD 8.6–9.5 26.7–24.2 Tomato; Durance and Wang, Sequence of combina- 2002 tion with hot-air drying MWVD 3.3–7.7 43–71 Cranberries; Sunjka et al., 2004 Microwave density and microwave mode MWFD 49.6 4.6 Sea cucumber Duan et al., 2010b

material is low (e.g., during the final drying stage). Consequently, MWFD is difficult to use in industrial applications due to the potential for plasma discharge problems, and it becomes necessary to control the process parameters (vacuum pressure and microwave power intensity) in order to avoid the appearance of this phenomenon. Duan et al. (2010b) suggested that pressure control in the range of 50–100 Pa, combined with microwave power decreasing in time, would ensure that no corona discharge occurred during microwave-assisted freeze-drying. It is important to design a chamber with as small a localized concentration of electromagnetic field as possible (Duan et al., 2010a). The lowest reported electric field intensities for dielectric breakdown in vapor and food have been 20 kV m 1 and 90 kV m 1,respectively (Cui, 2004). 304 9 Microwave-Assisted Drying of Foods – Equipment, Process and Product Quality

9.4.5 Changes in Dielectric Properties

As moisture is removed during the course of microwave-assisted drying, the dielectric constant e0 and the dielectric loss factor e00 normally decrease gradually, which can be seen in the descending parts of the curves of Fig. 9.17. However, during microwave-assisted freeze-drying, when most ice crystals have been removed, a sharp increase in e0 and e00 was observed by Regier and Schubert (2005) (cf. Fig. 9.17). This occurs because most of the water in the material is frozen during the initial period of the drying process, and this results in low dielectric properties (Wang et al., 2011a). In samples having the same moisture content, both,

Fig. 9.17 Variation of dielectric properties with time for restructured potato slices with different salt contents during MWFD. Reproduced with permission from Wang et al. (2011a). 9.4 Microwave-Assisted Drying Process 305 e0 and e00 are higher in the presence of salt than in its absence, which indicates a stronger microwave absorption capacity in the salty material.

9.4.6 Quality Changes in Food during Microwave-Assisted Drying

Following the water loss and temperature rises that occur during their microwave- assisted drying, food materials undergo complicated changes in their physico- chemical properties, such as color, shape, texture, volatile compounds, retention of nutrients and bioactive ingredients, as well as sensory quality. Generally, the color and nutrient compounds are degraded to some degree, while shrinkage or expansion in shape or structure also occur, depending on the type of material being dried. The shrinkage of seeds, grains, fruits and vegetables has been reported extensively. Shrinkage is related to moisture content because, when moisture is removed from the tissues, a pressure imbalance is created between the inside and outside of the tissue, generating compressive stresses that lead to shrinkage. Factors affecting shrinkage include pre-drying (initial moisture content), osmotic dehydration, sample thickness, microwave power, the final moisture content, and the type of microwave-assisted drying method used. It has been reported that osmotic pretreatment improved volume retention from 20% to 50% for strawber- ries, and from approximately 20% to 60% for apples (based on fresh volume), compared to samples subjected to microwave-assisted vacuum drying without osmotic pre-treatment (Changrue et al., 2008). Beetroot cubes were dehydrated by convective drying in hot air at 60 C and also by a combination of convective pre- drying (until the moisture contents were 1.6, 0.6 or 0.27 kg kg 1 d.b.) and microwave-assisted vacuum-finish drying at 240, 360, or 480 W. The use of microwaves at the end of the drying period significantly reduced shrinkage compared to the purely convective method. Increasing the microwave wattage and decreasing the time of pre-drying also improved the quality of beetroot cubes dried using the combined method (Figiel, 2010). Figiel (2009) also investigated the changes in the relative volume of garlic of various sizes, and found that drying to a moisture content of about X ¼ 0.6 was accompanied by a decrease in the relative volume of the material. Beyond this value no further change in the volume of garlic slices was observed, although for whole- and half-cloves a marked increase in volume was apparent. The shrinkage was greater at 240 W microwave power and moisture content over X ¼ 0.6, compared to samples dried at 480 and 720 W that had similar values of relative volume. Slices dried at 240 W until X ¼ 0.11 had the smallest relative volume of all samples dried with the vacuum microwave method. However, their relative volume was still greater than that of slices dried using the convective method. In contrast to the commonly encountered shrinkage in materials of plant origin, some gel materials – such as restructured fish muscle, imitation cheese, and restructured blue honeysuckle snacks containing starch – also display expansion. Figure 9.18 shows the typical expansion process of cheese containing starch during 306 9 Microwave-Assisted Drying of Foods – Equipment, Process and Product Quality

Fig. 9.18 Dynamics of temperature increase separate the regions of heating up, moisture (), weight loss (&)anddegreeof loss and expansion, and moisture loss and expansion (Î) during microwave heating of burning. Reproduced with permission from imitation cheese samples (fat : starch Arimi et al. (2008). ratio ¼ 0). Dashed vertical lines with arrows

microwave heating and drying. Three major effects were observed to occur: a temperature increase; expansion; and weight loss. The time course of the above processes may be divided into three distinct regions: heating up; moisture loss and expansion; and moisture loss and burning (Arimi et al., 2008). Typically, food products dried by microwave-assisted processes can display three types of structural change:

Shrinkage with porous structures such as most fruits, vegetables, and mush- rooms. A nonuniform pillow-shaped expansion with open structure, such as potato cubes produced by microwave-enhancedspoutedbeddrying,restructured blue honeysuckle snacks containing starch produced by MWVD, and microwave-expanded imitation cheese (Arimi et al., 2008; Yan, 2011; Liu et al., 2009). A uniform expansion with porous structures such as restructured fish muscle (Wang et al., 2013).

Both, shrinkage and expansion indicate damage of the original structure of the material, which can be undesirable if the dried food is to be consumed after rehydration. For a dried snack, however, expansion may be considered as a desirable indication of good crispness of the product. Several groups have correlated the volume or size changes of food materials during microwave-assisted drying with the moisture content (Tab. 9.3). 9.4 Microwave-Assisted Drying Process 307

Tab. 9.3 Some correlations for the change of volume of food materials during microwave- assisted drying; (see volume and chapter notation for symbols).

Sample Model Reference

CD- Beetroots V=Vi ¼ aX þ b Figiel, 2010 MWVD 2 MWVD Garlic V=Vi ¼ a bX þ cX Figiel, 2009 MWVD Button V=Vi ¼ a þ bXðÞ=Xi Giri and Prasad, 2006 mushrooms

MWCD Imitation ½ðÞV Vi =Vi 100 ¼ 70:72 0:012 Arimi et al., 2010 cheese exp½ 17:9X iP¼3 i MWVD Carrot d=di ¼ aiX ðradial shrinkageÞ Nahimana and Zhang, ¼ i 1 2011

In terms of end-product quality, foods processed by microwave-assisted drying often rank between air-dried and freeze-dried products (Orsat et al., 2007). Compared to convective drying, products dried via microwave-assisted processes have a reduced shrinkage, a faster rehydration with a higher capacity, a fresher-like color, greater retention of color or aroma, a higher nutritional value, and a softer or crispier texture. Moreover, microwave drying provides an effective bactericidal action in the dried product (Xu et al., 2004). Quality is further improved when vacuum is used, as the thermal and oxidative stresses are reduced. In some cases, MWVD can yield dehydrated products with characteristics that are similar to or even better than those dried by freeze-drying; for example, MWVD potatoes retain more fresh-like color than freeze-dried potatoes. Typically, the use of MWVD results in crispy and porous textures rather than the spongy and soft texture of freeze-dried foods. The impact of each microwave-assisted drying system on product quality is specific and heavily dependent on the drying conditions used. For example, microwave-assisted drying can conserve volatile aroma compounds at a high level when drying materials of plant origin, and can also produce good aroma and flavor in dehydrated fish products due to the development of aromatic aldehydes, including 2-methylpropanal, 3-methyl-butanal, and furfural (Wang et al., 2013). Rehydrated MWVD-restructured fish was inferior to products produced by conventional drying methods in terms of sensory evaluation, whereas rehydrated MWVD shrimp, carrot slices and russet potatoes were equal to or better than rehydrated freeze-dried samples. Microwave vacuum- dried cranberries showed a softer texture and were less tough than cranberries processed by microwave-assisted convective drying. MWCD cranberries were better appreciated by a sensory panel than was the MWVD product, although berries that had undergone both microwave drying methods were less desirable than hot-air-dried cranberries (Sunjka et al., 2004). The application of excessive microwave energy frequently leads to serious charring within the materials due to thermal runaway (as noted above). Charring is also closely 308 9 Microwave-Assisted Drying of Foods – Equipment, Process and Product Quality

related to the material’s properties, with heterogeneous, square, thick and shrinkable materials being susceptible to the development of charring spots due to their nonuniform dielectric properties (Wang et al., 2013; Holtz et al., 2010; Ahrne et al., 2007). By using optimal microwave-assisted drying conditions, the high-quality attributes desired by consumers and the excellent functional attributes of the dehydrated products can be provided and relatively easily maintained.

9.5 Microwave-Assisted Drying Process Control and Optimal Operation

9.5.1 Factors Controlling Microwave-Assisted Drying Processes

In order to achieve high-quality products with specific usage and sensory quality and also to save energy, a reasonable microwave-assisted drying process control is required, but this is a complicated and systematic task. The factors to be considered are shown schematically in Fig. 9.19, and have been discussed in Section 9.4. According to the scheme shown in Fig. 9.19, a relatively optimal drying process can be obtained by seeking a suitable dryer and drying method for a specific food product, as well as by selecting appropriate operating parameters.

9.5.2 Optimal Operation Strategy

Generally speaking, it is expected that multiple objectives are accomplished during the drying process, such as the least energy consumption and costs as well as the highest product quality. The quality characteristics and performance measures of the drying systems usually are described by several response variables that are greatly affected by many factors (see Fig. 9.19), and their interactions. Some of these variables should be maximized, while others should be minimized. In many cases, these responses are in competition one with another; that is, to improve one response may have an opposite effect on another response, further complicating the situation. Numerical optimization techniques, graphical optimization techni- ques and constrained optimization techniques have all been successfully used for process optimization (Shi et al., 2008). For numerical optimization techniques, either an overall desirability function or a synthetic evaluation index should be established. Desirability is an objective function that ranges from zero (least desirable) to one (most desirable) (Giri and Prasad, 2007b). Numerical optimization identifies a point that maximizes the desirability function or the synthetic evaluation index. Reported examples refer to optimization of the drying of edamames by a combination of hot air and vacuum microwave (Hu et al., 2006), optimization of the drying conditions for vacuum-assisted microwave drying of 9.5 Microwave-Assisted Drying Process Control and Optimal Operation 309

Microwave-assisted drying factors

Drying time Energy consumption Product quality

Operating Material Drying method parameters characteristics

Combination type Temperature Species Combination sequence Microwave emission mode Variety Microwave power Maturity Pressure level State Air velocity, Air humidity Composition Load and conversion point Blanching Osmotic treatment Shape Size Thickness

Process optimum

Optimal drying conditions

Scale-up

Fig. 9.19 Controlling factors of microwave-assisted drying processes, and process intensification strategy. green peas (Pisum sativum L.) (Chauhan and Srivastava, 2009), and optimization of the process parameters for the microwave vacuum drying of apple slices, using the response surface method (Han et al., 2010). In the graphical optimization technique, contour diagrams of the different response variables are superimposed and the optimal conditions are determined within the superposition region. Process parameters for the microwave freeze-drying of instant vegetable soup were optimized using this technique by Wang et al. (2010b). Sharma and Prasad (2006b) optimized the process parameters for the micro- wave-convective drying of garlic cloves using a constrained optimization technique (Tukey’s multicomparison pair-wise test), in an effort to identify the best compromise between quality and drying time. The results revealed that a 310 9 Microwave-Assisted Drying of Foods – Equipment, Process and Product Quality

microwave power of 40 W, an air temperature of 70 C and an air velocity of 1.0 m s 1 produced dehydrated garlic cloves of good quality and involved a low specific energy consumption during the drying process.

9.6 Concluding Remarks

Microwave-assisted drying provides unique opportunities for the development of advanced food-drying technologies. The main advantage of combining microwaves with other drying methods is to sharply reduce the drying time required for a product and, as long as the product temperature is controlled, then newly developed drying methods can be used to improve product quality. Unfortunately, most reports on technologies of microwave-related combination drying have been based on laboratory-scale systems, and there remains a need for further studies to bridge the gap between laboratory-based research and industrial applications. Currently, a large number of processing choices exists for the microwave-assisted dehydration of foods. Typically, each of these drying systems is product-specificin terms of end product quality and the functionality of its components, and consequently the choice of processing method will depend heavily on the intended use of the product and its market value. Yet, any given process can be detrimental to a particular product, with perhaps significant reductions in the retention of bioactive ingredients and/or their activities. Moreover, the mechanisms of product degradation and bioactivity are related to the processing conditions – principally of time and temperature and their control, with temperature-controlled microwave drying having been proven as the best processing method. In regard of materials, most studies on microwave-assisted drying have focused on agricultural products such as potatoes, carrots, and apples, although some have involved fabricated/formulated products. The main advantages of fabricated products, as noted by Gebhardt (1996), include reproducibility, uniformity and a lack of defects when compared to heterogeneous materials such as raw potato. It is clear that Life Cycle Assessments (LCAs) should be made of the various microwave- assisted drying techniques available, in an effort to determine a definitive selection strategy for dryers and drying processes. In addition, as concerns relating to climate change and global warming begin to affect dryer operations, the need for LCA analyses will become even greater in the future.

Additional Notation Used in Chapter 9

c speed of light m s 1

dp penetration depth m E electric field strength V m 1 9.6 Concluding Remarks 311

f frequency s 1 ¼ – MR moisture ratio, MR (X Xeq)/ (Xi Xeq) P microwave power density W m 3

Pabs absorbed microwave power W PMW microwave power W V volume m3 X moisture content kg water (kg dry matter) 1

Greek Letters e 1 0 permittivity of free space Fm e0 dielectric constant – e00 dielectric loss factor – e relative complex permittivity – l wavelength m

Subscripts i initial f final 0 freeze space

Abbreviations

CD convective drying d.b. dry basis FD freeze-drying LCA life cycle assessment MW microwaves MWCD microwave-assisted convective drying MWFD microwave-assisted freeze-drying MWSD microwave-assisted spouted bed drying MWVD microwave-assisted vacuum-drying SEC specific energy consumption w.b. wet basis 312 9 Microwave-Assisted Drying of Foods – Equipment, Process and Product Quality

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10 Infrared Drying German Efremov

10.1 Introduction

Drying (dehydration) is, by definition, the removal of moisture (as either water or other volatile compounds) from solids, continuous sheets, solutions and pastes, with the purpose of obtaining a solid product that is sufficiently low in water content. Drying is the most diverse among engineering unit operations, and also the most energy-consuming. The consumption of energy for drying of different materials depends on the type of drying process used. The common modes of heat transfer for drying are:

Convection from a hot gas in contact with the material or evaporation in vacuum. Conduction from a hot, solid surface in contact with the material. Radiation from a hot gas or surface. Heat generation within the material by dielectric, radiofrequency, or microwave heating.

Drying heat may be supplied by using electric or natural gas-fired radiators, or volumetrically in the microwave or radiofrequency range by placing the wet solid in radiation fields. By exposing an object to infrared (IR) radiation (wavelength of 0.78 to 1000 mm), the thermal energy generated can effectively be absorbed by materials. Infrared radiation can be felt as radiant heat when standing in front of a fire. As radiant heat flux can be adjusted over a wide range of wavelengths it is possible to obtain high drying rates for different surface-wet materials. In direct-heat dryers, the transfer of heat by convection from hot gases to the wet material is often supplemented by thermal radiation from hot surfaces that surround the material. This radiant-heat contribution is usually minor, and ignored. The use of thermal radiation as the major source of heat is a proven technology for the drying of certain films, sheets, textiles and coatings (Devahastin, 2000). During the 1950s A.V. Lykow and coworkers reported the results of their theoretical and experimental studies of infrared drying (Ginzburg, 1969).

Modern Drying Technology Volume 5: Process Intensification, First Edition. Edited by Evangelos Tsotsas and Arun S. Mujumdar. Ó 2014 Wiley-VCH Verlag GmbH & Co. KGaA. Published 2014 by Wiley-VCH Verlag GmbH & Co. KGaA. 318 10 Infrared Drying

Infrared drying provides significant advantages over conventional drying, including reduced heating time, uniform surface heating, reduced quality losses, versatile, simple and compact equipment, and significant energy savings. However, the most popular applications involve use of combined drying. Consequently, IR drying can be coupled effectively with vacuum and heat pump heating to permit the removal of evaporated moisture. Moreover, IR dryers may be used either continuously or intermittently to improve product quality. The aim of this chapter is to evaluate existing knowledge in the area of IR heating, to provide insight for the relationship between product properties and drying processes, and to present an up-to-date view on further research. The chapter also encompasses applications of IR heating in drying operations. Many industrial dryers can be modified to accommodate IR heaters; indeed, combined convective, microwave and IR dryers have been shown to be very attractive in this respect. Other combinations can also be used for drying.

10.2 Radiation Heat Transfer

10.2.1 General Principles

Radiation heat transfer refers to electromagnetic radiation, mainly in the IR range. From a practical viewpoint, the sole effect is thermal in this range – that is, heating of the receiving body by radiation from the emitting body. Above the absolute temperature of zero Kelvin (273 C), all substances emit electromagnetic radiation with the speed of light in a vacuum (c ¼ 2.998 108 ms 1). The frequency of the wave, f, is related to the wavelength, l, f ¼ cl; ð10:1Þ and varies as a function of the body temperature. The energy transmitted by the wave depends on its frequency. In contrast to conduction and convection, heat transfer by radiation does not require the presence of a material medium. Electromagnetic radiation is created by oscillatory electric charges (electrons and ions), and the frequency of oscillation determines the wavelength and type of radiation emitted. Radio, television waves and microwaves have long wavelengths, whereas ultraviolet, X-rays and gamma rays exist at short wavelengths. In between these is a range of wavelengths known as the optical thermal spectrum, with IR and visible light. Thermal radiation is, therefore, of the same nature as any other type of electromagnetic radiation, but has the strongest heating effect. The wavelength range of thermal radiation is 0.1 to 100 mm; IR radiation falls in this spectrum and is conventionally classified as near-infrared (NIR; 0.75–3.00 mm), medium infrared (MIR; 3.00–25 mm), and far infrared (FIR; 25–100 mm) (see Sandu, 1986). As shown in the spectrum of radiation (Fig. 10.1), the electromagnetic wavelength, l, depends on the manner in which it is generated. Heat transfer can 10.2 Radiation Heat Transfer 319

Fig. 10.1 Spectrum of radiation. be viewed as the propagation of electromagnetic waves, consisting of electric and magnetic fields that oscillate at right-angles to each other and to the direction in which the radiation travels. So, radiation heat transfer takes place by the transfer of electromagnetic waves from a hot, opaque surface through a nonabsorbing gas or vacuum to the material being dried. When radiant electromagnetic energy impinges on the surface of a body, it may induce changes in the electronic, vibrational and rotational states of atoms and molecules. The type of mechanism leading to energy absorption is determined by the wavelength range of the incident radiative energy and can be categorized as follows:

Changes in the electronic state corresponding to the wavelength range 0.2 to 0.7 mm (ultraviolet and visible rays). Changes in the vibrational state corresponding to wavelength range 2.5 to 1000 mm. Changes in the rotational state corresponding to wavelengths above 1000 mm (microwaves) (Decareau, 1985).

The broad region occupied by infrared extends from 0.76 mm (i.e., just beyond the red end of the visible end of the spectrum) to 400 mm. However, the radiation used for process heating occurs between wavelengths of 1 and 5 mm in order to obtain adequate source temperatures. The radiation emitted by a body can be determined if the temperature and nature of its surface (emissivity) are known.

10.2.2 Reflection, Absorption, and Transmission

A freely propagating electromagnetic wave transports the same amount of energy per second. When the wave enters a solid the increase in oscillations of the electrons causes the energy density to increase, so that the wave travels more slowly. An electric field is set up by the oscillating electrons, and this causes a part 320 10 Infrared Drying

Q Qρ n

Qα

Qτ

Fig. 10.2 Interaction of thermal radiation with material (incident energy flow rate, Q, absorbed energy flow rate, Qa, reflected energy flow rate, Qr, and energy flow rate transmitted through the material, Qt).

of the electromagnetic field to be reflected. Materials made only of atoms with tightly bound electrons absorb very little energy and are good heat insulators. In general, however, radiation is absorbed, reflected or transmitted by solid materials. By referring the absorbed energy flow rate, Qa, which is the energy converted to heat, the energy flow rate reflected from the surface of the material, Qr, and the energy flow rate transmitted through the material, Qt, to the incident energy flow rate, Q, the absorptivity, a ¼ Qa/Q, reflectivity, r ¼ Qr/Q, and transmissivity, t ¼ Qt/Q, can be defined (Fig. 10.2). It is, by definition a þ r þ t ¼ 1: ð10:2Þ Materials may be classified based on their absorptivity, reflectivity and transmis- sivity. A black body absorbs 100% of the radiation and does not transmit or reflect any (t ¼ 0, r ¼ 0, a ¼ 1). A body that transmits all radiation is termed “transparent,” and is characterized by t ¼ 1, r ¼ 0, a ¼ 0. A body that does not allow any radiation to be transmitted through it is termed “opaque,” and is characterized by t ¼ 0, r þ a ¼ 1. Whilst most solids are opaque, liquids and some solids such as rock salt or glass, on the other hand, have a high transmissivity and so are transparent to radiation. Only the absorbed portion of the incident energy causes heating. A black body is a body that emits the maximal possible amount of thermal radiation at a given temperature. The emissive power of a black body is a sole function of its temperature. Figure 10.1 shows that the peak wavelength emitted by a black body decreases as the temperature of the emitting body increases. The reflection may be either regular (also termed specular) or diffuse (scattering), which depends on the surface finish of the material. In the former case, the angle of incidence of the radiation is equal to the angle of reflection due to highly polished or smooth surface. When the surface has roughnesses larger than the wavelength, radiation is reflected diffusely in all directions. Generally, solid bodies absorb all of the radiation in a very narrow layer near the surface, and this is a very important consideration when modeling the heat-transfer process since, mathema- tically, this concept transforms a term within the energy balance into a boundary 10.2 Radiation Heat Transfer 321 condition (Mujumdar, 2007). Infrared heat transmission was used incidentally in the past accompanying other types of heat transfer during dehydration, but IR dryers are now designed to utilize radiant heat as the primary source.

10.2.3 Infrared Spectrum

Visible light is characterized by the fact that the radiation can be directed, focused and controlled by mirrors and lenses, and that prisms and gratings can be used for dispensing it into a spectrum. Ordinary sources of radiation in the visible spectrum (optical spectrum), such as tungsten filament lamps, fluorescent lamps and flames, consist of a very great number of molecules which have electric charges that oscillate independently of each other, producing a range of frequencies. The electromagnetic energy that is emitted from the surface of a heated body consists of a continuous spectrum of frequencies extending over a wide range. The spectral distribution and the amount of energy radiated depend on the temperature of the emitting surface. Measurements show that, for a given temperature, there is a definite frequency at which the radiated power is maximal (peak wavelength). The frequency of the peak is found to vary in direct proportion to the absolute temperature (Fig. 10.1). At room temperature, for example, the peak occurs in the far IR region of the spectrum, and there is no perceptible visible radiation emitted. However, at higher temperatures the maximal power is radiated at correspondingly higher frequencies, and at about 800 K a body begins to glow visibly. The properties of monochromatic radiation are termed “spectral,” and of polychromatic radiation “total” (Sandu, 1986). Hydrogen has relatively simple spectra, as the hydrogen atom consists of an electron and a proton. The electron has only 0.0005 the mass of the proton, and is able to inhabit only certain levels around the proton. To move from one level to another, the electron needs to gain or lose an amount of energy, called a quantum. For radio waves, a quantum is about 0.000 004 eV, for IR a quantum is about 0.004 eV, and for X-rays and c-rays it is about 40 000 eV. Electromagnetic radiation can be considered as stream of photons, which are massless particles traveling in a wave-like pattern. Movement between the lowest levels produces a photon of far-ultraviolet radiation, movement between the next lowest levels produces visible light and near-ultraviolet radiation, whilst movement between the middle levels produces IR radiation. The total amount of radiation emitted by a body per unit area and time is called the total emissive power, E, and this depends on the temperature and the surface characteristics of the body. This energy is emitted from a surface in all directions and at all wavelengths. A black body has, by definition, the maximal possible emissive power, Eb, which depends only on temperature. The emissivity of any body, e, is then defined as the ratio of its total emissive power E to that of a black body at the same temperature,

e ¼ E=Eb: ð10:3Þ 322 10 Infrared Drying

For a black body, the monochromatic emissive power (emissive power at a certain wavelength) is expressed by Planck’s law of radiation:

C1 E ;l ¼ ; ð10:4Þ b C l5 exp 2 1 l T ¼ 16 2 ¼ 2 where C1 3.745 10 Wm and C2 1.439 10 Km. fi e ¼ ’ The monochromatic emissivity of any body is de ned as l El/Eb,l. Kirchhoff s law states that, under thermodynamic equilibrium (which requires all surfaces to be at the same temperature), the monochromatic absorptivity and emissivity of any body are equal. Equation 10.4 has a maximum that is related to the temperature by Wien’s displacement law:

3 lmaxT ¼ 2:898 10 ðmKÞ: ð10:5Þ Planck’s law of radiation is displayed graphically in Fig. 10.3 for different temperatures and shows that, with increasing wavelength, the spectral fi intensity of the black body, Eb,l,is rst increased, then reaches a maximum, and finally decreases. The increase of temperature results in an appreciable

increase of Eb,l, and in a shift of the maximum (peak wavelength) to shorter wavelengths according to Wien’s displacement law (dashed line in Fig. 10.3). The black body radiates energy at every wavelength, so that the curves approach the abscissa of Fig. 10.3 asymptotically. At 5000 K (not depicted) the peak wavelength is about 5 10 7 m, which is in the visible light region, in the yellow-green section.

Fig. 10.3 Planck’s law of radiation. 10.3 Classification, Research, and Applications of Radiation Drying 323

Tab. 10.1 Total emissivity of various surfaces.

Material Temperature (C) Emissivity, e

Nominal or oxidized Polished

Aluminum 100 0.15 0.05 Brass 100 0.60 0.09 Steel 100 0.75 0.11 Stainless steel 100 0.85 0.17 Building brick 20 0.88–0.93 — Fireclay brick 500–1000 0.8–0.9 — Carbon black 50–1000 0.96 — Water 0–100 0.95–0.96 —

Equation 10.4 may be integrated over all wavelengths to obtain the total emissive power of the black body as a function of temperature to 1ð 4 Eb ¼ Eb;ldl ¼ sT ; ð10:6Þ 0 where s ¼ 5.67 10 8 Wm 2 K 4 is the Stefan–Boltzmann constant. According to Eq. 10.6 (Stefan–Boltzmann law), the total energy radiated by a black body is proportional to the fourth power of the absolute temperature. A gray body is defined as having the same emissivity over the entire wavelength spectrum. Thus, Kirchhoff’s law may be applied to gray bodies independently of their temperature. The emissive power of a gray body is given by

4 Eg ¼ esT : ð10:7Þ As the emissivity of real surfaces depends on the wavelength, average values over the relevant wavelength range are used in calculations. The average emissivity of various surfaces is listed in Tab. 10.1.

10.3 Classification, Research, and Applications of Radiation Drying

10.3.1 Classification

The various sources of electromagnetic radiation that can be used for drying include wavelengths ranging from the microwave to the visible spectrum (Fig. 10.4). Visible wavelengths are about 0.4 mm (violet light) to 0.7 mm (red light). The sun, the surface of which approximates a black body with temperature 5700 K, radiates most strongly in this range. Infrared wavelengths are longer than visible 324 10 Infrared Drying

Fig. 10.4 Spectrum of radiation for drying.

and up to about 100 mm. Thermal radiation is in the wavelength range of 0.1 to 1000 mm, and encompasses mainly the range of IR radiation. Infrared radiation can be further classified into the near-, mid- and far-infrared regions, as discussed above (see also Krishnamurthy et al. (2008)). In general, NIR waves are thermal and advantageous for use in drying, but MIR and FIR waves are less hot. For example, these wavelengths are used for the remote control units of televisions. Typically, the warmer an object, the more IR radiation it will emit. Drying heat may be supplied by using electric or natural gas-fired radiators, or volumetrically by placing the wet solid in dielectric fields in the microwave or radiofrequency range. As radiant heat flux can be adjusted locally over a wide range, it is possible to obtain high drying rates for surface-wet materials. Convection (gas flow) or vacuum operation is needed to remove the evaporated moisture. Infrared drying (4–8 mm wavelength) has found important applications in some niche markets, such as the drying of textiles, thick lumber, coated papers or printed sheets and films (Mujumdar, 2007). It is important to note that IR drying heat is absorbed at the surface, and that the application of continuous IR dryers is more common than that of batch IR dryers. In contrast to IR drying, dielectric drying involves the low-frequency, long- wavelength end of the electromagnetic spectrum (Fig. 10.1), where radio waves and microwaves reside. With electrically nonconducting materials, heat is not absorbed at the surface but rather is generated throughout the material, thereby reducing the importance of heat conduction within the material. This is making such drying unique, and also enables the rate of energy dissipation in the material to be controlled over a wide range. Other advantages over conventional drying methods include:

Efficiency of energy usage, because the energy dissipation occurs mainly in the moisture rather than in the solid material. Operation at low temperatures, thus avoiding high material surface tempera- tures. More rapid drying.

Dielectric drying is particularly useful for preheating materials and for removing the final traces of internal moisture to speed up drying at the tail end of the falling rate period. The main disadvantages of microwave drying include high capital and 10.3 Classification, Research, and Applications of Radiation Drying 325 operating (energy) costs. Only about 50% of the line power is converted into the electromagnetic field, and only a part of that is actually absorbed by the drying solid. Radiofrequency (RF) drying takes place at frequencies between 1 and 150 MHz (Seader et al., 2010). Although most materials are poor conductors of 50–60 Hz current, the impedance falls dramatically in the RF region. Such radiation can be used to heat the solid volumetrically, thus reducing the internal resistance to heat transfer. In RF drying, the energy is absorbed selectively by the water molecules.

10.3.2 Solar Drying

Solar radiation barely penetrates beyond the skin of the material, which absorbs only a part of the incident radiation depending on its wavelength. The term “solar dryer” is used for direct solar dryers, and is also reserved for a large variety of convective dryers whereby the products are not exposed to sun but are dried indirectly by air heated by solar energy. The convection of heated air in solar dryers may be either natural or forced. As a comprehensive treatment of solar drying has been provided in Chapter 6, Volume 4 of Modern Drying Technology, only a few examples will be described at this point. Sun drying at village level is common in tropical regions. Both, Bansal and Garg (1987) and Ekechukwu and Norton (1999) have referred to the dehydration of foods by direct exposure to radiation from the sun. Important quantities of fruits, vegetables, grains and fish are dried by this method, with sun-dried tomatoes being a specialty product of increasing popularity. A large proportion of the raisins, and practically all of the dried apricots and figs produced in the world, are sun-dried. Field-level experiments on the solar drying of pineapple using a solar tunnel dryer were conducted by Bala et al. (2003) in Bangladesh. The dryer consisted of a transparent plastic-covered flat plate collector and a drying tunnel connected in series to supply hot air directly into the drying tunnel, using two fans operated by a solar module. This dryer had a loading capacity of 120–150 kg of pineapple, and satisfactory drying times could be attained in a total of eight drying runs. The pineapples being dried in the solar tunnel dryer were completely protected from rain, insects and dust. Proximate analyses indicated that the dried product had a good quality for human consumption. In Japan, Basunia and Abe (2001) conducted thin-layer solar-drying experiments with medium-grain rough rice. In this case, the range of average drying air temperature was 22–35 C, the relative humidities were between 34% and 58%, and the initial moisture content of the rice was about 37%, dry-basis. A mixed-mode natural convection solar grain dryer was used, and data relating to sample weight and dry and wet bulb temperatures of the drying air were recorded continuously from morning to evening for each test. Subsequently, a good-quality dried product was obtained. Thin-layer solar drying experiments were also conducted by Yaldiz et al. (2001) for Turkish Sultana grapes, using an indirect forced convection solar dryer consisting of a solar air heater and a drying cabinet. The air velocity was varied, 326 10 Infrared Drying

and drying curves were correlated empirically. Experiments with thin layers of Turkish apricots were performed by Togrul and Pehlivan (2002) in an indirect forced convection solar dryer that consisted of a solar air heater with a conical concentrator and a drying cabinet. Air was forced into the solar air heater by a blower, and hot air obtained there was passed over the apricots. Changes in the mass of apricots and principal drying parameters were recorded continuously.

10.3.3 Infrared Drying

Infrared heating is widely used in industry (Mujumdar, 2007) for surface drying or for the dehydration of thin sheets such as textiles, paper, films, paints, coatings on ceramics, glass, and stone plates. In the automotive industry, IR drying for paint- on-metal applications and for paints on polyurethane-coated steel sheets has been highly successful. Another sector in which IR plays an important role is the pulp and paper industry. An example has been provided in Sweden, where this industrial sector imposes higher demands on energy consumption than any other. Consequently, a relatively new method that not only improved the paper quality but also achieved an energy-efficient drying involved the use of electrically operated IR radiation heating (Hannervall et al., 1992). Currently, IR drying – which employs the most efficient form of electromagnetic radiation for heat transfer – has been attracting interest in the agricultural industry because of its high thermal efficiency, fast heating response time, and direct absorbability by the material. In fact, this drying technology has appeared as a potential alternative to traditional heating methods for obtaining high-quality dried foods. Although the IR drying of thick porous materials has not yet been fully developed, some research groups have shown (mainly by means of experimental results) that one of the applications where long-wave IR heating is most efficient is in the dehydration of foods (Ginzburg, 1969; Kimura et al., 1992). In fact, IR radiation has been investigated for the drying of industrial grape byproducts (Celma et al., 2009b), seedless grapes (Caglar et al., 2009), olive husks (Celma et al., 2008), cashew kernels (Hebbar and Rastogi, 2001), herbs (Paakkonen et al., 1999), barley (Afzal et al., 1999), longan fruit (Nathakaranakule et al., 2010), potato (Afzal and Abe, 1999; Tan et al., 2001), carrot (Kocabiyik and Tezer, 2009), apple slices (Nowak and Lewicki, 2004; Togrul, 2005; Zhu and Pan, 2009), tomato byproducts (Celma et al., 2009a), shrimp (Fu and Lien, 1998), rough rice (Amaratunga et al., 2005; Abe and Afzal, 1997; Pan et al., 2008a; Khir et al., 2011), onion (Gabel et al., 2006; Jain and Pathare, 2004; Sharma et al., 2005), fresh and sugar-infused blueberries (Shi et al., 2008), persimmon (Kim, 1993), Ecklonia cava edible brown alga (Lee et al., 2010), saffron (Akhondi et al., 2011), pineapple (Tan et al., 2001), and red pepper (Koh et al., 1990; Nasiroglu and Kocabiyik, 2009). In general, these reports deal with the application of IR heating to the drying of different foods and analyses of the characteristics and performance of these particular IR dryers. In Japan, the food industry uses this type of heating for drying of, among others, seaweed, curry sauce, carrots and pumpkins (Kimura et al., 1992). 10.3 Classification, Research, and Applications of Radiation Drying 327

Fig. 10.5 Schematic diagram of laboratory-scale IR drying equipment. Reproduced with permission from Nasiroglu and Kocabiyik (2009).

Infrared drying provides significant potential advantages over conventional drying (Krishnamurthy et al., 2008), including a reduced heating time, uniform heating, reduced quality losses, an absence of solute migration in food material, versatile, simple and compact equipment, and significant energy savings. Infrared heating can be applied to various food processing operations, including drying, baking, roasting, blanching, pasteurization, and sterilization. In various reports, the effect of IR on food quality attributes has been discussed in the context of samples and process parameters. Applications of IR heating in food processing operations and future research potential are also reviewed. Infrared emitters with a maximal intensity wavelength in the range of 2.5 to 7 mm – that is, located in the MIR-to-FIR zone – have been proven suitable for the drying of agricultural products. A schematic diagram of the laboratory-scale IR drying equipment used by Nasiroglu and Kocabiyik (2009) is presented in Fig. 10.5. The drying chamber had a single-door opening at the front and was equipped with two incandescent lamps each with a power of 250 W. The output power of the lamps could be varied by regulating the voltage, and the air velocity could be adjusted by changing the fan revolution frequency. Before drying, samples were uniformly spread on the wire mesh tray in a thin layer. The sample mass and moisture loss were measured using a digital balance and recorded during drying. Infrared drying may be also classified by the type of emitter, whether an electric emitter (metal-sheath rods, quartz tubes, quartz lamps) or a gas-fired infrared (GIR) emitter (with ceramic-enclosed gas burners). Many previous reports have been primarily focused on the IR drying of agricultural materials using a conventional electric emitter (Abe and Afzal, 1997; Das et al., 2004; Jain and Pathare, 2004; Meeso et al., 2004; Sharma et al., 2005; Kumar et al., 2006). Investigations into paddy drying by using an IR emitter, in which heat was generated from natural gas (Amaratunga et al., 2005), involved the use of a gas-fired infrared (GIR) emitter that radiated wavelengths in the range of 2.5 to 3.0 mm. With 328 10 Infrared Drying

a maximum operating temperature of up to 900 C, the emitter was used to explore the drying kinetics of long-grain paddy, especially of premium grade rice (Khao Dok Mali 105 or Jasmine rice). The objective was to identify a model to describe the GIR radiation drying characteristics of this paddy type. Gas-fired IR drying of paddy was also considered by Laohavanich and Wongpichet (2008), who dried wet paddy under different process conditions by applying a pilot-scale experimental GIR dryer. The IR radiation was expressed in terms of peak wavelength of the IR emitter, and the initial moisture content of paddy was varied to study the drying behavior. Although, in general, sun drying and hot-air drying are the most common methods used for paddy drying, several groups (Ginzburg, 1969; Abe and Afzal, 1997) have aimed specifically at the IR drying of paddy. The effect of radiation intensity (0.125 to 0.5 Wcm 2) and slab thickness (2.5 to 10.5 mm) on the moisture diffusion coefficient of potatoes during FIR drying has been investigated by Afzal and Abe (1999). The study results showed that diffusivity was increased with increasing radiation intensity and slab thickness, whereas the activation energy for moisture desorption was decreased with increasing slab thickness, and this resulted in higher drying rates for slabs of greater thicknesses. Investigations into the IR radiation drying of high-moisture paddy and parboiled rice were continued by Das et al. (2004), who reported that both radiation intensity and bed depth had significant effects on drying performance and rice quality. X-ray microtomography (see Chapter 4, Volume 2, Modern Drying Technology), coupled with image analysis, represents a nondestructive technique that allows an entire sample to be scanned, providing information such as pore size distribution and total pore volume, without the need for serial cuts (as in the case of scanning electron microscopy). The use of X-ray microtomography was demonstrated when investigating the effects of FIR-assisted drying on the microstructure of banana (Leonard et al., 2008). The main advantages of IR dryers include:

They employ an alternative source of energy, and have a higher rate of heating. They have a high-efficiency conversion of electrical energy into heat for electrical IR. They heat only the object, and not the surroundings; selective heating is possible; and they can easily be subdivided to create zones for uniform heating of the surface-wet product. They respond very quickly to changing process conditions, and have quick start- up and shut-down procedures. Due to their smaller size they save on floor space; they are also easy to control and have lower capital and installation costs.

The potential limitations of IR dryers include:

They are difficult to operate with temperature-sensitive materials. 10.3 Classification, Research, and Applications of Radiation Drying 329

They have a low penetration depth, and cause skin formation and blistering of some coatings. As IR is basically a surface phenomenon, it is difficult to dry heavier coatings. Typically, they cannot be scaled-up in a straightforward manner; laboratory and pilot trials are needed to confirm any design changes.

10.3.4 Catalytic Infrared Drying

Recently, a new method of flameless catalytic infrared (CIR) drying has been developed, in which energy is generated by catalyzing natural gas or propane conversion with a platinum catalyst. Relatively few reports are available regarding CIR dehydration. Catalytic combustion occurs by allowing fuel (natural gas or propane) to enter the back of the air-tight heater pan, and then dispersing it through a catalyst pad at the face. At the same time, the air diffuses through the catalyst pad from the front, causing fuel oxidation. Catalytic combustion occurs without a flame, which can damage the product at a temperature below the ignition temperature. The electromagnetic radiant energy from CIR has peak wavelengths in the range of MIR to FIR radiation (3–7 mm) and is used for drying. The peak wavelengths match reasonably well with the three absorption peaks of liquid water, which could result in a rapid moisture removal. Because the CIR directly converts natural gas or propane to radiant energy, almost without losses, it is more energy efficient than the typical IR emitters which use electricity. A comparison of the drying of high-solids onion with CIR heating and forced air convection (FAC) heating was provided by Gabel et al. (2006). For this, sliced onions were dehydrated at 60, 70 and 80 C by means of FAC heating or CIR heating, with and without air recirculation. An onion sample (250 g for CIR, 150 g for FAC) of intact slices was arranged in a single layer on the drying tray at a loading of 2.5 kg m 2. In general, CIR both with and without air recirculation, had higher maximum drying rates, shorter drying times, and greater drying constants than FAC at moisture contents greater than 50% (d.b.). The CIR laboratory dryer arrangement used by Gabel et al. (2006) is shown in Fig. 10.6. This consisted of a drying chamber with a CIR emitter (Catalytic Drying Technologies, Independence, Kansas, USA) mounted from the top of the chamber. The sample was placed in an aluminum wave-guide (48 30 cm2, upper rim; 42 22 cm2, base perimeter), used to achieve a uniform heating of the entire product and placed on a drying tray. A balance (0.1 g accuracy) was placed beneath the drying tray to measure the product weight during drying. An exhaust fan was located on the top of the drying chamber for ventilation. Two fans (Fig. 10.6a) mounted on each side of the dryer were used to recirculate part of the warm air in the drying chamber. These fans pulled air from the top of the drying chamber and fed it back into the chamber through slits running the entire length of the sides of the drying chamber. The CIR emitter was preheated by an embedded electric 330 10 Infrared Drying

Fig. 10.6 (a) Catalytic infrared dryer. 1, 10, onion sample; 11, balance; naturalgas;2,gasflowcontrol;3,air (b) Construction of catalytic infrared exhaust; 4, blower; 5, CIR emitter; 6, emitter. Reproduced with permission from recirculation of warm air; 7, computer Gabel et al. (2006). controller; 8, wave guide; 9, thermocouple;

element. The natural gas intake was regulated by a gas control valve and a computer. Two thermocouples were used to measure the product temperature; each thermocouple was placed inside the flesh of an onion slice, which was placed within the innermost 10 cm2 of the onion bed. The average temperature was used as input to control the product temperature by turning the emitter on or off, which was achieved by opening and closing the gas supply valve of the emitter. The maximum operating temperature of the emitter was 500 C at wavelength 3.3 mm. The FAC dryer used in the tests was an electrically heated column dryer with a diameter of 33 cm. A fan blew heated air through an electric coil heater and then through the column. The product was placed in a circular mesh tray near the bottom of the column and suspended by wires to the balance that recorded product weight changes during drying. The electric heating coils were automatically switched on and off to maintain the set point temperature, in a similar fashion to the CIR dryer. The average air velocity for all of these tests was maintained at 0.5 m s 1. For CIR, the drying tray was placed in the preheated CIR drying chamber and the thermocouples were positioned as described above. The distance between the emitter and drying tray was 15 cm, with a maximal radiation intensity of 4752 W m 2. During CIR drying tests with air recirculation, both recirculation fans were switched on during the entire test. The targeted final moisture content of the dried onion was set at X ¼ 0.1 in this study. The CIR tests showed much higher drying rates than the FAC drying at X > 0.5 (Tab. 10.2). Below this moisture content, the rates of CIR drying were smaller than the FAC dryer rates, which indicated a lower moisture transfer in the onion slices. Raising the drying temperature caused an increase in the drying rate of both the CIR and FAC trials. In the CIR drying tests, the constant rate period was absent or very brief, because onions are hygroscopic and hygroscopic, and colloidal foods tend to quickly enter the falling rate period. Additionally, a more rapid surface drying by CIR may result 10.3 Classification, Research, and Applications of Radiation Drying 331

Tab. 10.2 Summary of onion drying results.

Drying condition Drying time to reach Max. drying rate X ¼ 0.5 (min) g(kg initial mass) 1 min 1

60 C CIR, recirculation 54.5 7.8 46.9 8.4a) 70 C CIR, recirculation 32.0 1.0a) 59.9 4.2a) 80 C CIR, recirculation 24.0 4.6a) 81.5 6.6a) 60 C CIR 47.3 7.5 47.5 10.5a) 70 C CIR 36.5 7.8a) 67.7 9.2a) 80 C CIR 18.0 2.8a) 83.4 4.6a) 60 C FAC 56.0 8.5 20.7 3.1 70 C FAC 51.0 1.4 23.9 3.9 80 C FAC 33.0 5.7 43.2 22.1 a) Significantly different when compared to FAC value at same temperature. in a quicker entrance to the falling rate period due to water diffusing very slowly to the surface of the onion. The dried onion q uality, measured as pungency degradation, was similar for bot h drying me thods a t C.60 Thand e 70col or analysis showed a be tter product olorc (whit er and less yellow ) at lower temp eratures for CIR a nd at higher temperat ure s for FAC. Th e browning that was vi sible may h ave be en cause d b y t he high er s urfac e heat fl u x of CIR heating and the longer proc ess times of FAC drying . Aer obic plate c ounts and coliform count s were not si gnifi cantly diff erent for the CIR- or FAC- drie d prod ucts, alt hough samples driedbyCIR hadsignifi cantly low er yeast and mold counts th an those dried by FAC. On basis of their results, Gab el et al . (2006) recommended the use of CIR during the early stages of onion drying. The Office of Industrial Technologies (OIT) in the Energy Efficiency and R enewa ble Energy Divis ion of th e US De partmentw.oit.doe of .govEnergy) (w w has proposed the use of a CIR drying system, employing IR energy in the 4–7 mm range (range of MIR to FIR radiation), to dry wood products. The catalytic combustion of natural gas creates IR energy without a direct flame. This is an advantage, because direct flame combustion may damage the product at a temperature below the ignition temperature. Reducing the moisture content with IR drying by transferring energy directly to the moisture instead of heating the air and surrounding metal structure of the drying unit requires less energy, reduces air emissions, and allows more product to be processed compared to conventional drying. When coupled with advances in remote sensing technol- ogy and programmable logic systems, wood fiber processors will have the ability to more carefully control the drying units. Subsequently, a large prototype unit has been constructed and tested with wood chips, sawdust, and a variety of agricultural products. As the CIR drying system was proven to dehydrate forest and agricultural products efficiently, the current focus is on the conveyance system for distributing the product evenly throughout the dryer. In addition to perfecting the conveyance system, the demonstration project will document the 332 10 Infrared Drying

energy, w as te and economic b enefi ts resulting from the constru ction and operation of a full -s cal e dehydration u ni t. One company specializing in the adaptation of flameless CIR drying technologies is, as noted above, the Catalytic Drying Technologies Co. (CDT, www .catalyticdrying.com). CDT is a division of the American Catalytic Industrial Group, Inc., a diversified family of companies aiming at the application of IR drying technologies to agricultural businesses and food processing. Since 1953, CDT has developed and patented IR drying systems for food, feed and fiber, and over 100 applications have been tested to demonstrate improvements in the quality, yield and throughput for dehydration, disinfestation, and thermal treatment. In the case of rice, a 2% increase in yield, a 10% increase in throughput and a 5% cost reduction were reported by using CIR drying instead of hot-air drying. The main advantages of CIR drying appear as higher yields, low capital and operating costs, higher throughputs, uniform drying efficiencies, a higher quality of end products, a lower environmental impact, and ease of control.

10.4 Types of Radiators

10.4.1 General Considerations

Sources of IR radiant heat at surface temperatures in the range of 600–2500 K are electrically heated metal-sheath rods, quartz tubes and quartz lamps, as well as ceramic-enclosed gas burners and catalytic IR burners. Infrared heaters are well- suited to a wide variety of applications, and designed to transmit large amounts of energy quickly and with a high efficiency. The particular wavelength of the IR radiation has a critical effect on the efficacy of the heating process, and IR heaters, which are optimally matched to the materials to be heated, can provide energy savings of up to 50% over other technologies. With a correct selection of the radiation source the IR energy requirements for drying can be reduced by up to 50%. The wavelength, surface temperature and spectral range for the main types of electric IR radiators are listed in Tab. 10.3. The terms “metal rod,”“ceramic,” “quartz or halogen” heating element describe the main components from which the heating element is made.

Tab. 10.3 Wavelengths of various IR radiators.

Radiator or heat source Peak wavelength (mm) Surface temperature (C) Spectral range

Metal rod radiator 2.8–4.4 400–750 MIR Infrared lamp >1.3 <1950 NIR Ceramic surface radiator 2.8–5.0 310–750 MIR Halogen lamp <1.4 <2200 NIR Quartz rod radiator 2.1 1100 MIR 10.4 Types of Radiators 333

Tab. 10.4 Spectral distribution of energy.

Temperature <2 mm2–4 mm >4 mm Typical emitter

600 C 2.2% 37.2% 60.6% Ceramic/Metal sheathed 900 C 13.0% 46.4% 40.6% Standard medium wave 1200 C 26.1% 46.9% 27.0% Carbon 1600 C 43.2% 40.1% 16.7% Fast response medium wave 2200 C 62.5% 28.7% 8.8% Short wave 2700 C 73.3% 21.0% 5.7% Halogen, NIR 3000 C 77.9% 17.6% 4.5% High-powered halogen, NIR

Depending on the temperature of the heating element, the IR emitter delivers distinctly different radiations at various wavelengths. It is important to select the correct emitter for the product, as the wavelength has a significant influence on the heating process. Short-wave radiation can penetrate deep into some solid materials and ensure a uniform and thorough heating, whereas medium-wave radiation is absorbed mostly in the outer surface and predominantly heats the surface.Medium-waveradiationisparticularlywellabsorbedbymanyplastics, glasses and especially water, and is converted directly into heat. The proportions of energy in characteristic wavelength ranges for typical emitters are listed in Tab. 10.4.

10.4.2 Electric Radiators

Electric IR radiators are most frequently used in drying for the whole IR-range of applications (including NIR). They have a large range of power output with an even emission over a large surface area (working width). Moreover, electric IR radiators are simply designed and clean installations. One of the most successful emitters is the quartz lamp, because it ensures high power densities, a maximum heat efficiency, flexible design parameters, and good controllability. Electric IR radiators have either round (see Fig. 10.5) or tube forms. Details of the geometric sizes and maximum outputs for different electric IR radiators of the Heraeus Noblelight Co. are provided in Tab. 10.5. The so-called nano XP emitters (Tab. 10.5) have extremely high linear power combined with a nanodiffusion layer as a primary reflector for thermal processing applications that need to operate under extreme process conditions (linear power up to 630 Wcm 1, temperature up to 1000 C). Short-wave emitters are single- or twin-tube NIR heaters that are well-suited for processes that require a quick start- up and shut-off at high temperatures, as they achieve their full operating efficiency within seconds. A high heating efficiency and heating times of 1.5 s make the carbon emitters the only medium-wave heaters with response times comparable to those of short-wave lamps. Suitable for all MIR applications, these heaters also allow the user to match temperatures to the optimal absorption wavelength for each 334 10 Infrared Drying

Tab. 10.5 Technical data of electric infrared heaters (Heraeus Noblelight, www.noblelight.net).

Infrared heater Format (mm) Heated length Total length Max. output range from-to (mm) from-to (mm) (W)

Nano XP Twin- 34 14 50–6000 195–6180 18 000 tube NIR Short wave Halogen 8 diameter 50–1000 110–1100 9000 NIR Single- 10 diameter 50–1000 110–1100 6000 tube IR Twin- 23 11 50–3000 200–3150 14 000 tube IR Carbon Round 19 diameter 100–1600 130–1730 6400 tube IR Twin- 33 15 100–3000 130–3130 10 500 tube IR Fast med- Twin- 34 14 500–3000 600–3150 12 000 ium wave tube IR Medium Twin- 18 860–1500 100–1600 2500 wave tube IR 22 10 80–2000 120–2200 4500 33 15 300–4400 350–4500 12 000 Long wave Twin- 22 10 80–2000 120–2200 3000 tube IR 33 15 300–4400 350–4500 7000

drying process, and are produced in lengths of up to 3 m. Taken together, these features eliminate the overheating and contamination of sensitive substrates. Fast-r esponse, med ium-wave emitters are t win-tube emitters that can b e su pplie d i n length s up to 6 m for us e in la rge modu les. Med ium-wa ve a n d long-waveemittersareeffi cient and stable f or a long operat ing lifetime, and the ir s pectrum is well su ited for he ating processes w ith most mate rials, wit h lengths up to 4.4 m. The question of which IR source is best suited to the process must be answered based on the spectral absorption of water. Figure 10.7 shows the spectrum (solid curves) of short-wave (NIR) and medium-wave (MIR) emitters operated with the same electric power, but at different heater temperatures in comparison to the absorption spectrum of water (dots). The MIR emitter and its range are more suitable for IR heat absorption in water. The correlation between thermal temperature and the color of an emitter is presented in Tab. 10.6 where, as expected, a red color corresponds to a low temperature, and yellow and white to higher temperatures. The so-called hybrid IR emitter combines a carbon emitter and a short-wave IR emitter (see Tab. 10.5) in a quartz-glass twin-tube con figuration. This combination allows the use of IR heat in processes that require both short-wave and medium- wave IR radiation, and the result is homogenous heating, matched to the application. In contrast to pure carbon emitters, the voltage can be flexibly adjusted. 10.4 Types of Radiators 335

Tab. 10.6 Correlation of thermal emitter temperature with color (Boston Electronics Corporation; www.boselec.com).

Temperature (C) Color (apparent)

400 Red, visible in the dark 474 Red, visible in twilight 525 Red, visible in daylight 581 Red, visible in sunlight 700 Dark red 800 Dull cherry red 900 Cherry red 1000 Bright cherry red 1100 Orange red 1200 Orange yellow 1300 Yellow white 1400 White 1500 Bright white 1600 Dazzling white (bluish white)

Fig. 10.7 Comparison of water absorption and IR emission spectra.

This combined emitter has advantages in applications where different absorption spectra or depths of penetration are needed.

10.4.3 Gas-Heated IR Radiators

Gas-heated IR radiators are conventionally used for the long- and medium- wavelength IR ranges, but are also applicable for short wavelengths. A high flexibility of the gas-heated radiator design can be attained by employing relatively small individual burners (suited to variable working widths). Typical gas-heated 336 10 Infrared Drying

Fig. 10.8 Gas-fired IR source. Reproduced with permission from Mujumdar (2007).

radiators consist of a perforated plate (metal or ceramic), which is heated by gas flames from the one side, so that the plate raises its temperature and emits radiant energy (Fig. 10.8). The porosity of the plate in fluences the temperature of the emitting surface so as to ensure a safe process. The radiant ef ficiency of such radiators is typically about 60% (Mujumdar, 2007). The company GoGaS Goch GmbH (www.gogas.com) develops gas-burning technologies for industrial applications, with a variety of types of IR burners covering the complete range of wavelengths of IR gas technology (see Tab. 10.7):

Long wave burner: These are catalytic burners that provide safe and flameless heating at uniform low intensity. A wide range of materials can evenly absorb long-wave energy, which allows the product to be heated in small increments. Medium-wave burners: These are ceramic or metal fiber burners with either a sintered or knitted (woven) surface. Medium-wave IR burners are reliable in operation and provide homogeneous irradiation to areas of almost any length and width. A higher energy efficiency can be achieved by using a hot-air cushion between the burner and the material to be dried. Short-wave burners: The concept of such gas IR burners (called RADIMAX) was developed in 2005. They differ from conventional burners in that combustion does not take place at the burner surface, but rather inside an inert three- dimensional porous medium. This combustion technology emits very intense IR radiation for applications in industrial drying and heating.

In gas catalytic IR emitters (Boone, 2005), unmixed fuel enters the back of an airtight heater pan and is dispersed through a catalyst pad which is preheated by a low-power electric element (Fig. 10.6b). Oxygen from the surrounding air diffuses through the catalyst pad from the front side of the heater. When the fuel gas and oxygen meet in the presence of the heated catalyst, an oxidation reaction occurs, generating heat. Once the catalytic reaction begins, it is self-sustaining and the electric preheater can be switched off. The surface temperature of these emitters is between 230 C and 480 C, a temperature range that places these emitters in the 10.5 Interaction between Matter and Infrared Radiation 337

Tab. 10.7 Parameters of various types of infrared burners (Company GoGaS).

Type Catalytic Ceramic Metal burner, Metal Porous burner burner fiber woven sur- burner, fiber burner face sintered sur- (RADIMAX) face

Range Long wave Medium Medium Medium Short wave wave wave wave Wavelength 3.3–5 2.4 2.2 2.2 1.7 (mm) Max. burner 600 950 1050 1050 1450 temperature (C) Max. thermal 30 120 200 250 1000 load (kW m 2) Dimensions 600 600a) 140 375 140 375a) 150 200 150 200 (mm) a) Special sizes on request. medium- and long-wave categories of IR. As the maximum surface temperature is below the auto-ignition temperature of the fuel, there is no flame (Boone, 2005). Another company offering gas-heated IR radiators is Thermo Catalytic Systems Ltd ww( w.thermocatalytic.com). Reported applications concern the accelerated curing of coatings and of heat-sensitive lacquers applied onto wood and wood- based products. Thermocatalytic heaters provide an economical source of heat, which is safe even in the presence of flammable solvent liquids and vapors. The basic types of emitter are designed to operate within defined limits of temperature determined by their construction and materials.

10.5 Interaction between Matter and Infrared Radiation

10.5.1 General Relationships

The design and modeling of any process always requires a profound knowledge of the materials. Specifically for IR drying, this fact may be the clue to accomplish a safe and efficient process, because the radiation properties of both the radiator and the material to be dried must be matched in order to obtain most efficient results (Mujumdar, 2007). Infrared radiant heat transfer to the surface of material is much like convective heat transfer. When the radiant energy reaches the material being dried, it is absorbed at the surface, from which it is transferred into the interior by conduction. Parts of thermal radiation may be reflected from the surface of material. 338 10 Infrared Drying

The absorptivity, a,reflectivity, r, and transmissivity, t, are the key radiation properties of a material. The magnitudes of a, r, and t depend on the type of material, its thickness and its surface finish, and also on the wavelength of the radiation. Absorption characteristics are defined in a similar manner to emission characteristics (see Section 10.2). The absorptivity for a nonblack body is the ratio of the nonblack absorption to the black absorption at the same surface temperature. Only the absorbed portion causes heating; however, as very few bodies behave as black bodies a more realistic assumption would be to treat those as gray bodies, which have the same absorptivity over the entire wavelength spectrum. Heat exchange by radiation between two black bodies at different temperatures may be obtained using the Stefan–Boltzmann law (Eq. 10.6) and is expressed by ¼ sð 4 4Þ; ð : Þ Qr AiFij T i T j 10 8 fi where Fij is the shape factor between surfaces i and j.Byde nition, this geometric factor takes into account the part of the total radiation emitted by the surface i that is intercepted by surface j. To calculate theoretically this factor is rather complicated, but for most common geometries various charts and formulas are available (Mujumdar, 2007). A useful equation, known as the reciprocity theorem, relates the shape factors and the areas of both surfaces:

AiFij ¼ AjFji: ð10:9Þ If the exchange of energy by radiation is between N bodies, the shape factors must follow the relation:

XN Fij ¼ 1: ð10:10Þ j¼1

The net radiation between two gray bodies is then given by the following equation: sð 4 4Þ ¼ T 2 T 1 : ð : Þ Qr r r 10 11 1 þ 1 þ 2 e1A1 A2F2;1 e2A2 According to Fig. 10.9, surface 1 in Eq. 10.11 is the surface of the material to be dried by absorbing radiation from the emitter 2, which is surface 2. Flat emitters (for example, gas-burning emitter, Fig. 10.9a) or lamp emitters (Fig. 10.9b) may be

(a) (b) S 2 2 D H H 1 1

Fig. 10.9 Arrangements of (a) flat and (b) tube emitters (2) and drying surfaces (1). 10.5 Interaction between Matter and Infrared Radiation 339 used. The distance H from emitter to material should be 1.5-fold more than the distance S from emitter to emitter. It should be noted that sometimes the electromagnetic radiation that impinges on a body may be attenuated inside the body by scattering along with absorption. Suppose that El(0) represents the spectral intensity of radiation impinging normally a layer of material where it is absorbed and scattered, so the intensity of monochromatic radiation is attenuated following the relation- ship (Lambert–Beer law): 2 3 ðz 4 5 ElðzÞ¼Elð0Þexp KlðzÞdz ; ð10:12Þ

0 where Kl is the extinction coefficient, and z is the distance. The extinction coefficient depends on temperature, pressure, composition, and the wavelength of the incident radiation.

10.5.2 Radiation Properties of Materials

The material to be dried by IR should have a low reflectivity in order to minimize the power required to heat it and, depending on the specificdrying process, a high or medium absorptivity. When drying thin materials or coatings, a high absorptivity of the material is usually better; however, when drying thick moist materials such as foodstuffs, a high transmissivity is preferable in order to avoid extremely intensive heating and thermal damage of the surface. It is important to point out that if the absorptivity of a material is low, its transmissivity is high, and vice versa. Properties such as the absorptivity and transmissivity of moist materials are not frequently encountered in the literature. In addition to the dependency with wavelength and thickness, they also depend on the water content. Figure 10.10 shows the IR absorption spectrum characteristic of water films with thicknesses of 2 and 10 mm. The spectral absorption of water films increases with the film thicknesses; generally, many fully wet materials will have their minimum absorptivity at those wavelengths where water has its maximum transmissivity, highlighting the important role that water plays in radiation absorption. The variation of absorptivity or transmissivity of moist materials with wavelength is difficult to estimate without experimental data. For many materials, transmissiv- ity is higher at lower wavelengths (Ginzburg, 1969). As an example, Fig. 10.11 shows experimental data on the total absorptivity of opaque building materials, depending on the temperature of the source (Mujumdar, 2007). Foodstuffs are complex mixtures of different large biochemical molecules and polymers, inorganic salts, and water (Sandu, 1986), and the absorption bands of each of these constituents are not the same. The transmissivity decreases with an increase in layer thickness, whereas absorptivity increases. An approximate way of representing experimental 340 10 Infrared Drying

Fig. 10.10 Spectral absorption characteristic of water films with thicknesses 2 and 10 mm. Illustration courtesy of Heraeus Noblelight Ltd.

Fig. 10.11 Absorptivity of some materials as a function of temperature. 1, white fireproof clay; 2, aluminum; 3, wood; 4, cork; 5, asbestos; 6, porcelain; 7, concrete; 8, graphite; 9, roofing silvers. Reproduced with permission from Mujumdar (2007). 10.5 Interaction between Matter and Infrared Radiation 341

Tab. 10.8 Transmissivity of selected materials.

Product Spectral peak, l (mm) Thickness (mm) t or tl (%)

Cigarette paper (raw) 1.075 — 73.0 3.220 — 48.0 6.150 — 20.0 Paper, thick (moist wall paper) 1.190 — 18.0 3.750 — 10.0 6.150 — 7.0 Wool cloth, dry 1.190 — 14.3 3.220 — 12.3 7.750 — 2.0 Sand 1.190 3.00 4.8 6.150 3.00 2.0 Wood 1.190 1.50 4.3 6.150 1.50 0.7

transmissivity data as a function of thickness is presented by Ginzburg (1969). Table 10.8 shows the experimental data for the variation of transmissivity of different materials commonly dried, together with thickness and peak wave- length (Mujumdar, 2007). In order to illustrate how properties of the emitter and the material to be dried should be matched during IR drying, Fig. 10.7 shows the radiant energy spectra of short-wave (2200 C) and medium-wave (900 C) IR emitters operated with the same electric power, together with the spectral absorptivity of water (Sandu, 1986). Consideration of the latter is essential for fast and good drying. The wavelength with the highest absorbency of water is a little above 3 mm,andthisregionisservedbymedium-waveradiatorswithalmost double energy as by short-wave radiators. Hence, medium-wave IR emitters are more suitable for IR heat absorption in water. Despite this fact, short-wave radiators are frequently used in practice, the main reasons being the lower initial costs and an easier technical realization. Short-wave radiators are readily available in different sizes and with the desired electrical tension. It should be remembered, however, that in the process of IR heating the transmission and absorption wavebands are usually sufficiently broad to utilize the side-bands of the emitted radiation in addition to the peak wavelength. As drying proceeds, the material being dried suffers a change in its radiation properties, increasing its reflectivity and, consequently, lowering its absorptivity at low water contents. It is then possible to change adequately the temperature of the emitter in order to improve the absorption of radiation during drying. If the effective thermal conductivity of the material is low, the surface temperature may rise to an undesirable value, particularly if high-temperature IR sources are used. Reports on the IR drying of foodstuffs were summarized in Section 10.3.3. For example, Krishnamurthy et al. (2008) reviewed aspects of IR heating and presented a theoretical basis for the IR heat processing of food materials and the interaction 342 10 Infrared Drying

of IR radiation with food components. Sandu (1986) highlighted the main advantages of IR drying of foods as the versatility of IR heating, a simplicity of the required equipment, easy accommodation of IR heating in convective, conductive, and microwave dryers, fast transient response, and also significant energy savings. The application of IR radiation in a stepwise manner by slowly increasing the power, with short cooling periods between the different power levels, resulted in less color degradation of bioproducts than with intermittent IR heating at the same power level (Chua and Chou, 2005). Reductions in overall color change of 37.6% and 18.1% were obtained for potato and carrot, respectively. Grinchik et al. (2003) considered the modeling of nonisothermal mass transfer in a porous system including periodic irradiation interaction. The long-term IR heat treatment may darken the color of onion due to browning (Gabel et al., 2006). Other important IR applications in the food area (not specifically in drying) are the cooking of soybeans, cereal grains, cocoa beans and nuts, precooking rice, bacon, and barley grains for ready-to-eat products, braising meat, and frying. An automatic IR system for the selective heating of foods, which could have a major impact on sterilization and pasteurization processes, has been investigated by Irudayaraj and Jun (2000).

10.6 Kinetics of Infrared Drying

The modeling of drying kinetics has been discussed in several previous volumes of Modern Drying Technology, especially in Volume 1. Nevertheless, the kinetics of IR drying of foods is almost always treated empirically in the primary literature, usually by regression on experimental data. An IR dryer was developed and used on onion slices at infrared power levels of 300 to 500 W, air temperatures of 35 to 45 C and air velocities of 1.0 to 1.5 m s 1 by Sharma et al. (2005). Afzal et al. (1999) investigated the IR drying of barley, whereas Das et al. (2004) evaluated the performance of a batch, vibration-aided IR dryer for high-moisture paddy. Midilli and Kucuk (2003) studied the thin-layer forced and natural solar drying of shelled and unshelled pistachio samples. Experiments of indirect solar drying of apricots with forced convection were evaluated by Togrul and Pehlivan (2002). Research on paddy drying by using an IR emitter, in which heat was generated from natural gas, was carried out by Amaratunga et al. (2005). Empirical equations have been used to correlate measured IR drying kinetics in all of these cases. Similar research on paddy has been conducted by Laohavanich and Wongpichet (2008), who used a pilot-scale gas-fired IR dryer in batch mode. The drying chamber had a length of 1.2 m, a width of 0.9 m, and a height of 1.2 m. A slide-up door allowed insertion and removal of the sample. The gas-fired IR emitter was installed on the top panel, and the radiated power could be controlled by using a natural gas flow regulator. Burned gas and moist air in the chamber flowed out through a vent hole, which was fixed at the top of the chamber. The emitter temperature was measured by thermocouples at the perforated plate of the emitter. 10.6 Kinetics of Infrared Drying 343

In the drying chamber, a paddy sample was placed on a wire mesh drying tray (0.9 m 0.45 m), to which a vibrating unit had been attached for mixing the paddy during the drying process. The distance between the radiator and the paddy sample was set constant at 40 cm in order to maintain a uniform radiant distribution throughout the experiments. The vibrating frequency was set at 450 cycles per minute and the amplitude of the vibration was approximately 10 cm in the vertical direction. Laohavanich and Wongpichet (2008) expressed the moisture ratio (MR), X X MR ¼ eq ; ð10:13Þ X 0 X eq where X0, X, and Xeq are the initial, current, and equilibrium moisture content, respectively, by means of the so-called Page equation to MR ¼ expðÞk tn ; ð10:14Þ where t is the elapsed time of drying, and k and n are fitting parameters. The parameters k and n were correlated to the IR peak wavelength and to the initial moisture content of the product, X0. The obvious disadvantage of Eq. 10.14, and of any other empirical equation, is that the values of k and n are valid for a specified set of experimental data, which means for a given particulate material and the range of operating parameters used in experiments. Extrapolation to other materials and drying conditions may lead to significant errors. The accuracy of predictions also depends on the number of experimental points, as the regression to determine k and n rarely reflects the process trajectory. Another limitation of Eq. 10.14 is that at the start of drying (t ¼ 0), the reduced moisture content MR calculated from the right-hand side term is not necessarily equal to one, and therefore a significant discrepancy is noted between experiment and simulation in the drying kinetics. Some limitations of the empirical equations could be eliminated when using a semi-empirical approach, with the process model built on the theory of heat and mass transfer, but model parameters determined from experiments. A good example is the so-called modified quasi-stationary method (MQSM) that has been used to model drying kinetics of different materials (Efremov, 1999; Efremov and Kudra, 2005). The MQSM is resulting in the following equation for drying kinetics 1 MR ¼ ; ð10:15Þ 1 þ ðÞt=s n where t is the elapsed time of drying, s is a characteristic time, which is assumed to be constant for given process conditions, and n is a dimensionless exponent, which reflects the effect of the heating. Equation 10.15 is a new form of the generalized drying curve with the reduced time (t/s) and the index of intensity of heating (n). The semi-theoretical origin of Eq. 10.15 allows more accurate predictions to be made of drying kinetics with a minimum of experimental data, and also enables the preheating drying period to 344 10 Infrared Drying

be described. In addition, the extrapolation of limited experimental data is far more reliable, as Eq. 10.15 reflects the real variation of the material moisture content with drying time. If the index of intensity of heating, n, and the characteristic time, s, are already determined from existing data, then Eq. 10.14 can be rearranged to calculate the time of drying: 1=n t ¼ s MR 1 1 ð10:16Þ and the rate of moisture content change, N ¼dX/dt, which is proportional to the drying rate,

n1 n 2 t t N ¼ ns 1ðX X Þ 1 þ ; ð10:17Þ 0 eq s s for any moisture content. To validate the MQSM, the MR calculated from Eq. 10.15 was compared with the previously mentioned experimental data for wet paddy (KDML 105 variety) dried under different process conditions in a pilot-scale gas-fired IR dryer (Laohavanich and Wongpichet, 2008). The drying experiments were conducted

using 0.5 kg of wet paddy samples at three initial moisture contents, X0, of 0.22, 0.27, and 0.37. Infrared radiation was applied at a specific maximum emitted l ¼ m wavelength of max 2.47 m, which was close to the absorption peak of moist materials in the IR ray range, corresponding to an emitter temperature of 900 C.Themoisturecontent,X, was obtained from the experiment at each testing condition for drying times, t,of0,0.5,1.0,1.5,2.0,3.0,5.0,7.0,and 10.0 min. These experiments were conducted at ambient inlet air temperatures of approximately 28–30 C. Table 10.9 compiles the calculated characteristic drying time, s, for the IR drying kinetics of paddy at different initial moisture contents, as depicted in Fig. 10.12. The value of n for MQSM was also determined from the experimental data. Because the heating conditions during these experiments were identical, a constant value of n ¼ 1.386 was obtained for all three drying curves. Figure 10.12a presents a comparison of the computed curves with experimental drying kinetics obtained for paddy dried at constant IR radiation intensity but at different initial moisture contents. The moderate spread of experimental points

Tab. 10.9 Characteristic drying time (s) for IR drying kinetics of paddy at different initial moisture content.

Parameter Symbol in Fig. 10.12 ~

Initial moisture content, X0 0.22 0.27 0.37 Characteristic time, s (min) 5.477 4.438 3.677 Coefficient of determination, R2 0.9993 0.9994 0.9993 10.7 Infrared Drying Combined with other Types of Drying 345

Fig. 10.12 Comparison of experimental and predicted moisture ratio (MR) at three initial moisture contents (see Tab. 10.2) for IR drying of paddy (Laohavanich and Wongpichet, 2008), evaluated by MQSM. (a) Separate curves; (b) Generalized drying curve.

around the respective computed lines indicates that Eq. 10.15 describes very well the kinetics of drying for material with preheating period finished at t ¼ 1.5 min. Figure 10.12b presents the drying curve MR ¼ f (t/s) generalized for three initial moisture contents, as specified in Tab. 10.9. A discernible but marginal deviation of experimental points from the generalized drying curve indicates that Eq. 10.16 can be used to compute the drying time of paddy, even beyond the range of the parameters tested in these experiments. The drying of paddy occurred mainly during the period of falling drying rate. The MR-values could be predicted with a high value of the coefficient of determination, R2 ¼ 0.9993, using the MQSM. In considering the prediction of drying kinetics for IR drying by MQSM, future studies should be aimed at quantifying the exponent, n, by experiments on drying the same material in different dryers, followed by the drying of different materials in the same dryer. In the next step, an attempt should be taken to determine the characteristic time, s, from material properties and the principles of radiation mass transfer.

10.7 Infrared Drying Combined with other Types of Drying

Infrared radiation is almost always present in heat-transfer processes, and can be considered as a surface treatment similar to other radiation technologies. Often, the IR source has a high temperature (500–3000 C), and heat transfer by convection is also taking place and cannot be ignored. As penetration of this type of radiation is poor, the heating effect of IR radiation has an impact almost only on the surface of the body, with heat transfer through the body proceeding by either conduction or convection. An overview of combinations of IR drying with other drying methods is provided in Tab. 10.10. 346 10 Infrared Drying

Tab. 10.10 Infrared drying combined with other types of drying.

Type of dryer Product Reference

IR þ Convective Polyvinyl alcohol Putranto et al., 2010 Herbs Paakkonen et al., 1999 Barley Afzal et al., 1999 Onion Sharma et al., 2005 Onion Kumar et al., 2006 Onion Gabel et al., 2006 Fruit leather Jaturonglumlert and Kiatsiriroat, 2010 IR þ Microwave Honey Hebbar et al., 2003 Carrot and garlic Baysal et al., 2003 Peach Wang and Sheng, 2006 Potato and carrot Chua and Chou, 2005 IR þ Freeze Banana Pan et al., 2008b Aloe vera Chakraborty et al., 2011a Tiger prawn Chakraborty et al., 2011b Sweet potato Lin et al., 2005 Strawberries Shih et al., 2008 IR þ Vacuum Banana Swasdisevi et al., 2009 Onion Mongpreneet et al., 2002 IR þ Fluidized bed Soybean grains Dondee et al., 2010 Carrot Honarvar and Mowla, 2012 IR þ Heat pump Banana, guava, potato Chua et al., 2001 Logan Nathakaranakule et al., 2010 IR þ Vibration Paddy Das et al., 2004 IR þ Superheated steam Banana Nimmol et al., 2007

10.7.1 IR and Convective Drying

As shown in Tab. 10.10, the most popular combination involves use of convection (hot air) and IR radiation. It is often reasonable to boost the drying capacity of an existing convective dryer for herbs (Paakkonen et al., 1999), sheets such as paper, fruit leather, or film of polyvinyl alcohol (Putranto et al., 2010) by using IR heating. When Kumar et al. (2006) considered the IR and hot-air drying of onion slices, the latter were dried under different processing conditions, namely slice thickness, inlet air temperature and air velocity. On exploring the IR and convective drying of onion slices, Sharma et al. (2005) observed that the drying time for onions increased with the increase in air velocity at all IR powers applied; however, the drying time was reduced with an increase in IR power. The application of combined IR radiation and conventional convective heating is considered to be more efficient than radiation or convective heating alone, as it provides a synergistic effect. Afzal et al. (1999) reported that during the combined convective and IR heating process of barley, the total energy required was reduced by about 1.56-, 2.38-, and 10.7 Infrared Drying Combined with other Types of Drying 347

2.45-fold as compared to convection drying alone at 40, 55, or 70 C, respectively. Hebbar et al. (2004) developed a continuous combined IR and convective dryer for vegetables, whereby the synergistic effect of IR and hot air led to a rapid heating of the materials, resulting in a higher rate of mass transfer. The evaporation of water took 48% less time and required 63% less energy for combined mode drying as compared to convective drying. Gabel et al. (2006) compared the drying and quality characteristics of sliced high-solids onions dried with catalytic infrared (CIR) heating and forced air convection (FAC) heating. CIR, both with and without air recirculation, showed higher maximal drying rates and shorter drying times than FAC at moisture contents greater than 50% (cf. Section 10.3.4). Jaturonglumlert and Kiatsiriroat (2010) considered heat and mass transfer for the combined convective and FIR drying of fruit leather, and found that the ratio between the heat and mass transfer coefficients for the combination technique could not be obtained from the heat-mass transfer analogy. Hence, modified correlations for predicting the ratio of heat and mass transfer coefficients in term of the heat transfer Nusselt number were developed.

10.7.2 IR and Microwave Drying

Baysal et al. (2003) studied the effects of microwave and IR drying on the quality of carrot and garlic, evaluated by instrumental and sensory analysis, and compared with the effects of conventional hot air. Rehydration, moisture content, water activity, particle density, bulk density, porosity and color values wereobtained,andthedryingrateversusfreemoisturecontentwascompared for the three different drying methods. When Wang and Sheng (2006) compared the far-IR with microwave drying of peach, the dehydration rate of the fruit was increased and energy consumption decreased with increasing microwave or IR power. The peach tissue experienced two falling rate periods when dried with microwaves or far-IR, with the transition at X ¼ 1.7. The same water loss consumed more energy at X < 1.7. Unlike microwave drying, an accelerating dehydration rate period occurred immediately after the start of far- IR drying (at about 8 < X < 9). The application of microwave (2450 MHz) and NIR radiation (IR lamps 1.0 kW, peak wavelength 1.1–1.2 mm) was explored for the thermal processing of honey, and the effects on the physico-chemical characteristics and microbiological quality of honey was studied by Hebbar et al. (2003). The conventional method involves the indirect heating of filtered honey at 60–65 C for 25–30 min in a tubular heat exchanger, followed by rapid cooling in order to protect the honey’s natural color, flavor, enzyme content and other biological substances. Microwave heating provided a rapid means of achieving the desired level of yeast reduction with reduced thermal damage (the reduction in moisture content was above 9% at power levels of 50, 70 and 100 W, when the samples were heated for 60 s). Infrared heating achieved the desired results in a relatively shorter period (3–4 min), thus offering advantages over the conventional method. The use of IR heating was 348 10 Infrared Drying

sufficient to obtain a commercially acceptable product, which met all the statutory requirements of quality. A comparative study between intermittent microwave and IR drying of bioproducts was conducted by Chua and Chou (2005). The application of IR radiation in a stepwise manner, by slowly increasing the power and inserting short cooling periods between the successive power levels, resulted in less color degradation. Reductions in overall color change of 37.6% and 18.1% were obtained for potato and carrot, respectively.

10.7.3 IR and Freeze-Drying

Sequential infrared and freeze-drying (SIRFD) as a new processing method was studied by Shih et al. (2008) for producing high-quality crispy fruit pieces at reduced cost. The drying characteristics of strawberry slices and the quality of the finished products were investigated. Samples predried by IR to X ¼ 0.3, 0.4 and 0.5 were then further freeze-dried to achieve final moisture content of about X ¼ 0.05. When the color, shrinkage, rehydration ratio and crispness of the finished products were measured, those prepared using SIRFD had more desirable properties than freeze-dried fruit without IR predehydration. Pan et al. (2008b) also studied the SIRFD process for banana slices. Samples predehydrated with 20% and 40% mass reduction by using 4000 W m 2 IR intensity were further subjected to freeze-drying for various times to determine the effect of predehydration on the drying rate. The quality characteristics of dried banana chips, including color, thickness shrinkage and crispness, were studied. The predehydration results showed that the drying rate of IR heating was significantly higher than the hot-air drying rate and increased with the increase in IR intensity. It has been concluded that SIRFD can be used for producing highly crispy banana chips; additional acid dipping improved the product color and reduced the required freeze-drying time. A combination of IR heating with freeze-drying was shown to reduce the processing time of sweet potatoes to less than half (Lin et al., 2005). A shorter wavelength resulted in rapid drying and thus reduced the drying time. Empirical equations were developed to describe the drying rate during the falling rate period. Chakraborty et al. (2011a, 2011b) studied the IR-assisted freeze-drying of aloe vera and of tiger prawn (Penaeus monodon).

10.7.4 IR with other Types of Drying

Combined far-IR radiation and vacuum drying has recently received more attention as the technique has proved effective in drying some types of fruits with the aim of producing fat-free fruit-based snacks. However, most studies are purely experi- mental, for example the drying of welsh onion by FIR radiation under vacuum conditions (Mongpreneet et al., 2002). Swasdisevi et al. (2009) considered a 10.7 Infrared Drying Combined with other Types of Drying 349 mathematical model, which allows prediction of the changes of moisture content and temperature of a model food material (banana) undergoing combined IR and vacuum drying. The predicted results were compared with experimental results at absolute pressures in the drying chamber of 5, 10, and 15 kPa and at controlled banana center temperatures of 50 C, 55 C, and 60 C. The model performed adequately in predicting the changes of moisture content. The simulated temperature at the center of a banana slice was higher than the experimental temperature during the first 50 min of drying, but agreed well afterwards. The drying of a cylindrical sample in a fluidized-bed dryer containing inert particles was studied by Honarvar and Mowla (2012). For this purpose, a pilot- scale fluidized-bed dryer was constructed in which two different heat sources – hot air and IR radiation – were applied, and pieces of carrot were chosen as test samples. The heat-transfer coefficient for cylindrical objects in a fluidized bed was also measured. The heat absorption coefficient for carrot was studied; this can be computed by dividing the heat absorbed by the carrot by the heat absorbed by an aqueous solution of black ink. Subsequently, a mathematical model was proposed based on mass and heat transfer phenomena within the drying sample, and the results obtained in this way were in favorable agreement with the experimental data. Dondee et al. (2010) considered NIR radiation combined with fluidized-bed drying to reduce the cracking and breakage of soybean grains, based on its advantages such as a decrease in the moisture gradient and stress development within the grain kernel, leading to a high-quality product. The physical qualities (i. e., cracking, breakage and color) were investigated together with the microstructure of soybean grains; protein solubility and urease activity were also determined. The parameters studied were NIR radiation powers of 4, 6 and 8 kW, air velocity of 4.5 m s 1, air temperature of 40 C, and a grain bed depth of 6 cm. The results showed that the drying rate increased in line with increases of NIR radiation power, while the moisture content was reduced from an initial value of 20% to final moisture contents of 13.5% (4 kW), 12.8% (6 kW) and 12.5% (8 kW). The cracking and breakage of soybean grains (4.4% and 5.3%, respectively) seemed negligible, while the protein solubility of the final product was within an acceptable range (77–82%) for feed meal. The reduction of residual urease varied in the range of 39 to 70%. These results confirmed that soybean grains can be efficiently treated by combined NIR radiation and fluidized-bed drying. A highly sophisticated multistage heat pump with PID control was used by Chua et al. (2001) to achieve sinusoidal and square-wave-type temperature variations of the drying air. However, it may be possible to achieve similar benefits in a batch dryer by an intermittent supply of energy using IR lamps that switch on and off at a desired frequency and amplitude of heat flux. Because of the thermal inertia of IR sources, this method of energy supply is limited to low-frequency variations of heat supply. Alternative options for intermittent heating are volumetric heat generation by using microwave or RF irradiation. Although this mode of heating allows higher frequencies of intermittent energy supply to be attained, the capital costs of microwave and RF equipment are generally higher than for IR equipment. 350 10 Infrared Drying

Fresh longan must be preserved in the form of a dried product to prevent damage by microorganisms. A new technique of using FIR radiation in combination with hot air and heat pump dryers for longan was proposed by Nathakaranakule et al. (2010), whereby longan with an initial moisture content of approximately 85% (w.b.) was dried to about 18% (w.b.) at a drying temperature of 55 C in the case of combined FIR and heat pump drying, and at 65 C in the case of combined FIR and hot-air drying with 80% recycled air. In both cases, the electric power supplied to the FIR rods was set at 250, 350, and 450 W. The experimental results showed that FIR, in combination with hot air and heat pump drying, increased the drying rate and reduced the drying time of longan. FIR can also help to create a more porous structure in dried longan, with porosity increasing in line with the power supplied to the FIR heaters. The formation of a porous structure results in a product with lower shrinkage, improved rehydration, lower hardness and lower toughness than samples dried in the absence of FIR; the FIR-dried longan also had a stronger red color and appeared darker. It is worth noting that the overall energy used for FIR-assisted drying processes decreased with the increase in the power supplied to the FIR heaters. A batch-type vibration-aided IR dryer was developed (Das et al., 2004) and the drying characteristics of three varieties of high-moisture paddy were studied using 0.8 kg of each at radiation intensities of 3100 and 4290 W m 2 and grain bed depths of 12 and 16 mm. The drying rate was found to depend on the level of radiation intensity, but not on the bed depth. The drying took place almost entirely during the falling rate period. Nimmol et al. (2007) studied a drying system which combined the already proven concept of low-pressure superheated steam drying (LPSSD) with FIR radiation for banana. The effects of various operating parameters (i.e., drying medium temperature and pressure) on the drying kinetics and heat transfer behavior of banana, as well as the energy consumption of the process, were investigated and discussed. Comparison was also made with similar sets of data obtained with combined FIR radiation and vacuum drying, and from a system using only LPSSD. The results showed that LPSSD-FIR and vacuum-FIR combinations required shorter drying times compared to LPSSD under all drying conditions. In terms of the specific energy consumption, the specific energy consumption of the vacuum pump was seen to be much higher than that of the FIR radiator or electric heater. The specific energy consumption of the LPSSD-FIR and vacuum-FIR combinations were also found to be lower than that of LPSSD, under all drying conditions. Based on the drying rates and specific energy consumption of all tested processes, LPSSD-FIR at 90 C and 7 kPa was suggested as the best combination. Datta and Ni (2002) discussed the application of combined IR, microwave, and hot-air heating to food materials. The development of a continuous drying apparatus equipped with FIR heaters, NIR heaters and hot-air blast can reduce the economic costs, drying time and operating temperature. However, the vegetable size should be restricted to no more than 5 mm in thickness to improve drying efficiency. Salagnac et al. (2004) considered the numerical modeling of heat and mass transfer in a porous medium during combined hot-air, IR and microwave drying. In this case, the material is submitted on its upper side to hot air flow, as well as to 10.8 Conclusions 351

IR and microwave radiation. The one-dimensional physical model describing heat and mass transfer allows access to the temperature, moisture content and pressure fields, while energy inputs induced by electromagnetic radiation were determined by the Lambert–Beer law. A comparison of experimental and simulated results on cellular concrete demonstrated the relevance of this model to studies on combined drying.

10.8 Conclusions

Although IR drying is a unique process, the application and understanding of IR heating in industrial processing is at present still in its infancy. It is evident that IR heating offers many advantages over convection heating, including greater energy efficiency, heat transfer rate, and heat flux that results in time saving as well as increased production line speed. The modeling of IR heat transfer is a difficult task, as two or even three heat-transfer modes may occur in a system at the same time. Some thermal properties of materials can also vary with changes in temperature, while further problems may arise because a material is seldom regular in shape, or homogeneous. However, certain assumptions and empirical relationships have allowed the development of practical models that describe drying processes with sufficient precision. Most empirical heat-transfer models applicable to drying are highly specific, and general analytical models are seldom encoun- tered. Infrared heating is attractive primarily for surface-heating applications. In order to achieve an energy optimum and efficient practical applicability of IR heating in industrial processing, combinations of IR heating with microwave and other common conductive and convective modes of heating hold great potential. Indeed, it is quite likely that the use of IR heating in industry will be augmented in the near future, especially in the area of drying. Today, the use of IR heating is expanding rapidly on a worldwide basis.

Additional Notation Used in Chapter 10

c speed of light m s 1 E emissive power (intensity) W m 2 F shape factor for radiation – f frequency s 1 K extinction coefficient m 1 ¼ – MR moisture ratio, MR (X Xeq)/(X0 Xeq) s characteristic drying time s 352 10 Infrared Drying

Greek Letters

a absorptivity – l wavelength m r reflectivity – t transmissivity –

Subscripts

a absorbed b black body g gray body l spectral r reflected t transmitted 0 initial

Abbreviations

CDT catalytic drying technologies CIR catalytic-infrared db dry basis FAC forced air convection FIR far-infrared GIR gas-fired infrared IR infrared LPSSD low-pressure superheated steam drying MIR middle infrared MQSM modified quasi-stationary method NIR near-infrared OIT Office of Industrial Technologies PID proportional integral derivative RF radiofrequency SIRFD sequential infrared and freeze-drying wb wet basis

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Index

a antioxidant properties, of Saskatoon absorption 293 berries 150–151 – decrease of power 293 apple juice, foam-mat drying 182–185 – of energy 283 apple tissue – infrared drying 319–321, 338–340 – air-drying time reduction 219 – ultrasonic system 240 – continuous ultrasound 221 acoustic coupling, liquid media 245 – drying curves 201, 220 acoustic energy 239 – freeze-drying of 149 acoustic field 244 – impact of PEF on 198 acoustic impedance, ultrasound waves – impedance measurement 201 199, 240 – micrographs 203 acoustic pressure 239, 258–259 – osmotic dehydration 206 – dependence on air velocity 261 – porosity levels 192 – variations of 260 – temperature in microwave drying 298–299 acoustic stress, contact ultrasound 218 – ultrasound treatment 200 aerodynamic valves, pulse combustors 34–35 – water distribution 215 AFD. see atmospheric freeze-drying Archimedes number agitation, levels in acoustic medium 258 – and fluidization behavior 126 air, for feed foaming 173 – spouted beds 94, 99–100 air circulation 238 Arrhenius behavior, effective diffusivity 264 air decontamination, ultrasound ascorbic acid content applications 248 – under LPSSD and VD 74–75 air drying – of potato chips 62 – contact ultrasound 218–221 – quality parameters 212 – pretreatment 206–208 atmospheric freeze-drying (AFD) 132, – processes 256 146–149, 237 – time reduction by ultrasound 219 – kinetics of 147 air heating, energy for 1 – sample composition dependence 271 air relative humidity, AFD 149 – ultrasound-assisted 153, 265 air spaces, plant tissue composition 192 atmospheric-pressure SSD 61–69 air temperature 2, 263–266 atomization – as function of drying rate 18 – foam production 165 – influence of drying rate, air velocity, and – in PCD 37 nozzle diameter 20–21 ––secondary 38 air velocity atomizer, flux of liquid phase behind 41 – AFD 149 attenuation – effect on ultrasonic drying 259 – intern 339 – influence on air temperature 20–21 – ultrasound waves 240 airborne ultrasound 216 average Nusselt number, of single nozzle 6

Modern Drying Technology Volume 5: Process Intensification, First Edition. Edited by Evangelos Tsotsas and Arun S. Mujumdar. Ó 2014 Wiley-VCH Verlag GmbH & Co. KGaA. Published 2014 by Wiley-VCH Verlag GmbH & Co. KGaA. 358j Index b – temperature modeling 298 – bacteria inactivation undergoing LPSSD 71, 73 – – – LPSSD 74, 77–78 undergoing VD 72 73 – – microwave drying 307 caseinate, foamed 169 170 – ultrasound waves 200 caseinate powder 166 banana chips, produced by LPSSD-FIR 76–82 catalytic combustion 329 – banana foams catalytic infrared (CIR) drying 329 332 – drying curves 185 cavitation 200 – – foaming additives 186 contact ultrasound 217 – bananas, evolution of moisture content liquid-solid systems drying 253 – – LPSSD-FIR 79 missing in gas-solid-systems 256 – – VACUUM-FIR 79 stable 241 – barriers, minimization of mass transfer transient 241 fi – 191–236 cell damage quanti cation 192 195 – bed pressure drop, spouted beds 98 frequency-conductivity spectra 194 beetroot cubes, quality changes 305 cell disintegration – bell pepper tissues, drying curves 220 air drying 207 – biological surfactants 182 freeze drying 210 – biomaterials, superheated steam-drying methods 194 – of 57–82 osmotic dehydration 205 black body, absorption of 320 cell disruption, ultrasound waves 200 boundary layer cell membranes, limiting mass transfer 191 – stripping of 38–39 cell permeabilization, plant tissue drying 193 – thickness reduction 256 cell rupture, measuring degree of 193 fl boundary layer effects, ultrasound 242 CFD. see computational uid dynamics 30 brix sucrose solution chamber. see combustion chamber; drying – acoustic pressure 258 chamber; pressurized chamber, vacuum – liquid-solid systems drying 253 chamber bubble methods, foam production 164–165 chamber pressure, in primary freeze-drying bubbling fluidized bed design 86–87 stage 136–137 bulk foams, drying of 163 chaotic ultrasonic field 247 bulk volume structure, of foam-dried cheese – – particles 166 cottage 172 173 burning, of products 302 – quality changes 306 cheese-making processes 243 c chewiness, relationship with diffusivity 271 calorimetric methods 251–252 chips capacity-controlled limiting case 120–121 – banana 76–82 – carbon dioxide. see CO2 potato 62 – – carbon emitter, electric radiatiators 333 SSD 61 64 – – b-carotene content, under LPSSD, VD, and tortilla 61 62 HD 74–75 CIR. see catalytic infrared drying b-carotene depletion, quality parameters 212 CO2, feed foaming 167 carrot cubes, optimal ultrasonic energy 261 coffee, freeze-drying 138, 154 carrots coffee powders, foam spray-drying 179 – effective diffusivity 268 cold trap, microwave freeze-drying 290 – effective moisture diffusivity 257 color degradation 342 – freeze-drying of 151 combined drying – – gas-solid system drying 256 convective, compared to contact – hot-air drying 207 drying 112 – – – impact of pulsed electric fields (PEF) 198 freeze-drying 150 154 – – mass transfer coefficients 262 infrared drying 318 – – – osmotic dehydration 205 quality characteristics 224 226 Index j359 combustion. see also pulse combustion CSTR. see continuous stirred tank reactor Ò – catalytic 329 cyclone, of Pulse-Dryer 52 – continuous, versus pulse combustion 29–30 cyclone dryer 114 combustion chamber 27 cylindrical bodies, nozzles for 14–16 combustor-dryer system, geometry of 30 complex permittivity, relative 283 d compression tests 198 damping coefficient, ultrasound waves 199 computational fluid dynamics (CFD) models, defoaming, ultrasound applications 248 PCD 42–48 dehydration 279 condensation methods, foam production 164 – osmotic 205–206, 217–218 condenser, exergy losses in 137 – solar drying 325 conduction, infrared drying 317 dehydration rate, infrared and microwave conduction-controlled limiting case 120–121 drying 347 conical spouted bed 87, 96 DEM-CFD. see discrete particle modeling – cylindrical 89 (DPM) ––characteristic dimensions of 96 density, of powders 181 constrained optimization techniques, design space method, freeze-drying 137–139 microwave drying 308 detergents, porous matarials as 175–180 contact heating, intensification by 112–126 dielectric breakdown 302–304 – general principle 112–114 dielectric constant – main effects and influences 114–121 – food materials 284 – modeling 121–126 – ice 295 contact ultrasound 216–217 dielectric drying, infrared drying 324 contact ultrasound-assisted freeze-drying dielectric heating, energy absorption 283 system 224 dielectric materials, volumetric heating 281 contact ultrasound effects, on mass dielectric properties transfer 204 – changes 304–305 continuous combustion, versus pulse – microwave-assisted drying 283–285 combustion 29–30 diffuser. see tailpipe continuous (Eulerian) CFD 107 diffusion, cell membranes 191 continuous freeze-drying 132, 154–155 diffusion coefficients, air drying 207 continuous stirred tank reactor (CSTR), diffusion model ProCell units 103–106 – mass load densities 262 convection – modelling air velocity 259 – hot gas 317 diffusivity. see also effective diffusivity – infrared drying 346–347 – for apple 248 – moisture removal 324 – sample composition dependence 268 convective drying dipolar relaxation, microwave heating 282 – compared to contact drying 112 direct-heat dryers, infrared drying 317 – equipment 286–287 discrete element method (DEM) 107 – hot-air 207 discrete (Lagrangian) consideration, of particle convective heat transfer, in PCD 42 phase 107 conveyor belt, contact ultrasound 218 discrete particle modeling (DPM), fluidized- core, frozen, shrinkage of 266 bed processes 107–111 cosolvents, organic, use in freeze-drying disintegration indices, cell damage 144–146 measurements 193–194 cottage cheese powder 172–173 disk-distributor spray-dryer 50 coupling, acoustic system and medium 245 dispersed foams, drying of 163 cranberries, drying energy efficency 301 dispersion methods, foam production 164 crispness, of banana chips 81 dissipation, of electrical energy 283 crispy product texture, microwave distilled water, contact ultrasound 218 drying 280–281 DPM. see discrete particle modeling cryoprotectants 210–211 dried cake, thickness of 134, 38–139 360j Index dried layer, thickness of 238 – kinetics 257 driving force, of PCD 38 – reduction 207 droplet size. see Sauter mean diameter (SMD) – relationship with power 299 droplets dual-gas-entrance spouted beds 91 – foamed 163, 165, 169, see also foam dyes, plant tissue drying 193 materials – moisture content 47 e drying chamber effective diffusivity 269 – pressure in 133 – atmospheric freeze drying 266, 271 –– primary freeze-drying stage 136–137 – drying temperature 264 – vibrating 216 – frozen samples 271 drying conditions – kinetics modeling 257 – effect on quality 307 – modelling air velocity 260 – saskatoon berries 289 – ultrasonic influence 268 drying curves – water 253 – apple tissue 201 effective moisture diffusivity – comparison of different methods – dependence on acoustic power 257 293–295 – modelling air velocity 260 – contact heating 116–121, 123 egg albumin – single-particle 122 – additive for banana foams 186 drying energy. see energy – as surfactant 182 drying equipment 285–292 egg white, foam-mat drying of 186 – convective 286–287 eggplant drying, kinetics modeling 257 – microwave assistance 279–316 electric emitter 327 – spouted bed 291–292 electric field effects, during freezing 209 drying intensification. see also specific drying electric pulses 195 processes/techniques electric radiators, infrared drying 333–335 – process variables 250–272 electrical conductivity response, cell damage – sample composition dependence 268 measurements 193 drying kinetics electrical consumption, microwave-assisted – dependence on acoustic power 257–258 drying 300–301 – equation 343 electroacoustic sensors, ultrasonic field – infrared 342 measurements 252 – temperature dependence 263 electrofreezing. 210 drying methods, comparison of different 294 electromagnetic radiation, infrared drying processes drying 318 – combined 216–217 electronic state, infrared energy – microwave-assisted 279, 292–308 absorption 319 – multistage 66 electropermeabilization, PEF drying rate pretreatment 205 – catalytic 330 electroporation – constant microwave power 297 – cell damage results 195 – energy consumption as function of – freeze drying 210 18–19, 21 emissive power 321–322 – fluidized-bed dryer 349 emitter types, infrared drying 327 – influence of nozzle diameter 22 energy – microwaves 293 – for air heating 1 drying stages 222 – required by freeze-drying 133 drying temperature. see temperature – specific drying energy 4 drying time energy absorption – calculation 344 – infrared 319 – carrots, dependence of acoustic power 256 – microwave-assisted drying 283 – infrared and convective drying 346 – by microwaves 282 Index j361 energy consumption 237. see also specific flapper-type valves, pulse combustors 28, energy consumption 33–34 – for different methods 302 flat products, drying by nozzle arrays 1 – in drying 279 flow cytometry, plant tissue drying 193 – as indicator for ultraound efficency 251 flow design, intensification of fluidized-bed – microwave-assisted drying 299–302 processes by 86–111 – nozzle array 1 flow patterns, spouted beds 88 – reduced 58, 69 fluidization, characteristic regions of 93 – ultrasonic drying measurements 264 fluidized-bed processes 216 energy conversion, in dielectric materials 283 – AFD 147, 149 energy dissipation, ultrasound effects 241 – discrete particle modeling 107–111 energy efficiency – drying curves 115–121 – infrared drying 324 – infrared drying 349 – moisture dependence 300 – intensification of 85–127 – pulsed mode microwaves 301 ––by apparatus and flow design 86–111 energy loss, during ultrasound freezing 223 ––by contact heating 112–126 energy reflected, acoustic impedance 240 – microwave systems 291 energy savings, optimzed infrared heaters 332 – spray zone of 107–111 energy source, ultrasound drying 243–250 foam-mat drying 163, 181–187 energy transformation, ultrasonic – of egg white 186 equipment 243 – foaming gas 182 enhanced quality. see quality improvement – freeze-drying 185–187 enthalpy of vaporization 3–4 foam spray-drying 163, 169–170 ethanol removal 266 – final product properties 172–181 European Process Intensification Center – microscopic images of powders 166, 179–180 (EUROPIC) 27 – milk powder 174–175 evaporating mass flux 2 – processing principles 167–172 evaporative freezing 132 – structures of foamed particles 165 excitation amplitude, ultrasound intensity 219 foamed materials exergetic analysis, of freeze-drying 133–139 – drying of 163–187 exergy losses, during primary freeze-drying – production methods 164–165 stage 136–138 – properties of 164–166 expansion, food quality changes 305–306 foamed particles, structure of 165 external cavitation avoidance 225 foaming, effect on spray parameters 169 foaming additives, dying of bananas 186 f foaming gas far-infrared radiation (FIR)-assisted LPSSD – foam-mat drying 182 (LPSSD-FIR) 76–82 – foam spray-drying 169, 173 feed foaming foaming system, gas-admixing 167–168 – effects on final product properties 172–181 foams, of high stability 182 – main methods 167 food industries, freeze-drying feedback temperature control 299 intensification 131–156 fiber-optical temperature measurements 295 food matrix, PEF performance 196 fiber-optical temperature sensors 286 food products, penetration depths 285 field intensities, pulsed electric field foods (PEF) 195 – combined heating dryers 350 field measurements, ultrasound drying 251 – dielectric constant 284 field strength distribution, local 196 – freeze-drying 151–152 fields, stationary 252 – microwave-assisted drying 279–316 final product quality, effect of drying – physicochemical changes 305 temperature on 182 – quality changes during drying 305–308 FIR. see far-infrared radiation – superheated steam-drying 57–82 flameless catalytic infrared (CIR) drying 329 forced air system, microwave dryer 286 362j Index freeze-drying – quality changes 305 – atmospheric (see atmospheric freeze-drying) – temperature in microwave drying 298 – chamber pressure in primary stage 136–137 gas-admixing foam spray-drying 170 – combined technologies 150–154 – moisture content after 171–172 – contact ultrasound 221–224 gas-admixing foaming system 167–168 – continuous 154–155 gas bubbles, foam production 164 – design space method 137–139 gas-desorption foam spray-drying 169–170 – electric field effects 209 gas-fired infrared emitter 327, 336 – energy required by 133 gas-heated infrared radiators 335–337 – equipment 290–291 gas jet, pulse combustion 27 – exergetic analysis (and optimization) gas-solid systems, ultrasonic drying 133–139 effects 256 – foam-mat 185–187 gas stream, heat transfer coefficient in 40 – infrared drying 348 gas velocity, spouted beds 86, 94 – intensification of 131–156 gas volume flow rate, spouted beds 94–95, 97 – maximum product temperature 138 gaseous emissions, from pulse – microwave-assisted 150–152 combustors 30 – during primary drying stage 134 gaseous media, ultrasound drying 242, 247 – pulsed electric field (PEF) 208–211 gel materials, quality changes 305 – reference temperature 134 ginseng, freeze-drying of 151 – ultrasound-assisted 152–153, 224 glass beads, microwave systems 292 – use of organic solvents and cosolvents gooseberry, comparison of LPSSD and 144–146 VD 74–75 – vacuum 139–146 graphical optimization techniques, microwave freeze-drying cycle, development of 137 drying 308 freeze thawing 194 gray body 323 freezing, as one step in process – net radiation 338 intensification 131–132 Groenewold experiments 114–121 freezing currents 209 – summary 125–126 freezing tolerance improvement 211 freezing variables 208 h frequency conductivity spectra 194 HAD. see hot air-drying frequency range hardness – infrared drying 317 – of banana chips 81 – microwaves 281 – relationship with diffusivity 269 – sound 239 heat absorption coefficient, fluidized-bed – ultrasound waves 240 dryer 349 frequency-tunable pulse combustors 35–36 heat flow rate, in contact drying 113 – rotary kiln with 48 heat generation, infrared drying 317 frozen foods, penetration depths 285 heat pump drying (HPD), shrimps 66–69 fructose distribution, quality parameters heat-sensitive products, combined drying 213–215 technologies for 150–154 fruit juice, foam-mat drying 182–185 heat transfer 319 fruit pulps, foam-mat drying 184 – dependent on relative nozzle distance 9 fruit tissues – in hole channel 13 – contact ultrasound 218 – in individual nozzles 8, 18 – drying 192 – infrared drying 317 fl – – fuel ow rate, effect on NOx emission 31 in nozzle arrangements 10 11 fuel jet, pulse combustion 27 – for perforated plates 14 – in ProCell units 100–107 g – pulse combustion 38–42 garlic – radiative 318–323 – drying energy efficency 301 – ultrasound effects 242 Index j363 heat transfer coefficient – intracellular and extracellular 208 – for circular area 5 – removal, dielectric changes 304 – dimensionless 119 ice nucleation, control of 140–143 – heater-to-bed 118 ice sublimation 143, 145, 148 – in nozzle array 11–12 IMC. see intermediate moisture content – in oscillating gas stream 40 imitation cheese, quality changes 306 – ratio to mass transfer coefficient 2 immersed tubes, for contact heating 112, 125 – and specific enrgy consumption 4 impedance difference – wall-to-bed 118 – gaseous media 247 heating – liquid media 245 – contact 112–126 – solid media 247–248 – dielectric 283–285 impedance measurement, apple tissue 201 – volumetric 280 impinging jet drying 1–24 heating shelf, temperature of 133, 137 – application 1–4 Helmholtz combustor 32 – nozzle field design 17–23 high-intensity ultrasound 199, 252–253 – nozzle fields 7–16 – process intensification 250 – Nusselt functions 16–17 high-pressure air source, microwave – single nozzle 4–7 systems 292 improved drying, mass transfer barriers high-stability foams 182 191–236 high temperatures, quality effects 237 in-line arrangement, of nozzle arrays 8–11 hole channels individual nozzles, heat transfer in 8, 18 – nozzle arrays 1, 3, 12–13 induction heating 126–127 – specific energy consumption 20 infrared drying 317–356 holes, in microwave oven 291 – benefits 327–328 honey quality, infrared and microwave – catalytic 329–332 drying 347 – combined with different methods 345–351 hot-air drying (HAD) – convective drying 346–347 – bacteria inactivation 77–78 – different types of drying 348–351 – b-carotene content 75 – equipment 327 – convective 207 – freeze-drying 348 – shrimps 65 – kinetics of 342–345 – sugar beet pulp 64 – limitations 328–329 – ultrasound-assisted 153 – microwave drying 347–348 – ultrasound waves 202 infrared heaters, optimzed 332 hot gas convection, infrared drying 317 infrared radiation, interaction with hot spots, danger of 281 matter 337–342 HPD. see heat pump drying infrared radiators, gas-heated 335–337 hybrid drying technologies 70–76 infrared regions 324 hybrid infrared emitter, electric infrared spectrum 321–323 radiatiators 333 infrared temperature measurements 295 hydration, during ultrasonic drying 256 intensification, of drying processes. see specific hydrodynamic mechanism, contact processes/techniques ultrasound 217 intensity zones, ultrasonic field hydrophones measurements 252 – acoustic pressure measurements 259 interaction, between ultrasound and – ultrasonic field measurements 252 sample 266–272 hypertonic solutions 253 interface microstirring, liquid-solid systems drying 253–254 i interfacial polarization effects, cell damage ice, dielectric constant 295 measurements 193 ice crystals intermediate moisture content (IMC) 66 – fracture by cavitation 222 intermittent heating 349 364j Index intermittent microwave-convective m drying 299 magnetic resonance imaging (MRI) intermittent ultrasound treatment 225 – cell damage measurements 195 – inversion phenomenon, SSD 59 60 – quality parameters 213–214 inversion temperature, SSD 60 magnetostrictive transducers 243 investment costs, PEF systems 196 maltodextrin powders 176 ion leakage, ultrasound treatment 200 – final properties 177 ionic conduction, microwave heating 282 – microscopic images 166, 179–180 mass flow rate, in contact drying 113 j mass flux, evaporating 2 – jet drying 1 24 mass load, dependence on ultrasonic energy 261 k mass transfer kinetic energy, disordered 282 – electroporation 205 kinetics – hydrodynamic mechanism 217 – drying modeling 257, 296 – in ProCell units 100–107 – – infrared drying 342 345 – pulse combustion 38–42 Kirchhoff’s law 322 – ultrasound effects 241–242 mass transfer barriers l – process-induced minimization 191–236 laboratory dryer system, CIR drying – reduction 228 – 329 330 mass transfer coefficient 269 Lambert-beer law 339 – atmospheric freeze drying 271 layer development, during freeze-drying 271 – diffusion model including resistance 262 lemon peel – kinetics modeling 257 – effective diffusivity 268 – ratio to heat transfer coefficient 2 – effective moisture diffusivity 257 mass transfer resistance, reduction by – gas-solid system drying 256 temperature 263 liquid atomization, in PCD 37 mass transport liquid media, ultrasound drying 245 – increase by ultrasound, low-temperature fl liquid phase, ux behind atomizer 41 drying 265 liquid-solid systems, ultrasonic effects 253 – liquid-solid systems drying 253 liquid spray, pulse combustion 27 material surface temperatures, infrared liquids drying 324 – quasi-adiabatic state 252 material temperature, microwave heating 283 – – vacuum freeze-drying of 139 146 matter, interaction with infrared local heat transfer, in nozzle arrangements 10 radiation 337–342 local Nusselt number 9 maximum product temperature, freeze- fi long wave burner, gas- red infrared drying 138 source 336 meat brining, contact ultrasound 217 “looseness” of powders 176 mechanical cell disintegration 194 loss factor mechanical effects – penetration depths 285 – on tissues 227 – temperature dependence 283 – ultrasound efficency measurements 252 – low-intensity ultrasound 199, 252 253 mechanical oscillations, ultrasound 239 low-pressure superheated steam drying mechanical valves, pulse combustors 32–34 – (LPSSD) 69 76, 350 mechanical waves, piezoelectric effect 244 – – compared to VD 70 76 medium, effect on acoustic waves 258 – – inactivation of microorganisms 74, 77 78 medium characteristics, ultrasonic – – quality improvement 76 82 intensity 258 low-temperature drying, mass trasnport medium-wave burners, gas-fired infrared increase by ultrasound 265 source 336 lyophilization. see freeze-drying melon cubes, contact ultrasound 218 Index j365 membrane pores, mass transfer – power density 283 processes 191 – pulsed mode 298 metal-sheath rods 332 – relative complex permittivity 283 metal-shielded thermocouples 295 milk, freeze-drying 154 metallic particles, inert, for contact milk powder, foam spray-drying 167, 174–175 heating 126–127 milk protein, as surfactant 182 microbial inactivation. see bacteria inactivation minimization, mass transfer barriers 191–236 microchannels, gas-solid system drying 256 modeling microjets – of AFD 147–148 – liquid-solid systems drying 254 – air velocity, microphone measurements 260 – ultrasound effects 241–242 – contact heating 121–126 microorganisms, inactivation under – discrete particle 107–111 LPSSD 74, 77–78 – of PCD 42–48 microphones modified quasi-stationary method (mqsm), – modelling air velocity 260 kinematic models 343 – ultrasonic field measurements 252 moisture content microstirring 250 – after infrared drying 330 microstreaming, water removal – apple tissues 254 improvement 203 – of droplets 47 microwave-assisted drying 292 – as function of spray-drying time 171–172 – basic principles 281–283 – normalized 122 – controlling factors 308–310, 309 – pork loin 255 – dryer systems 287–288 – reduction 237 – electrical consumption 300 – relationship with power 299 – equipment 285–292 moisture diffusivity – food quality changes 305–308 – effective 257, 260 – foods 279–316 – temperature dependence 263 – inert medium fluidized-bed dryer moisture loss 324 system 292 – microwave-assisted drying 293–295 – infrared drying 347–348 moisture ratio 220 – optimal operation 308–310 – drying time calculation 345 – process control and optimal operation – modeling in infrared drying 343 308–310 moisture removal – techniques 238, 280 – microwave heating 282 microwave convective drying (MWCD) 280, – rapid 280 286 moisture transfer 207 – air-drying systems 287 moisture uptake capacity, specific, of air 113 microwave freeze-drying (MWFD) 150–152, monochromatic emissive power 322 279, 290 monochromatic radiation 321 microwave generator 286–287 MQSM. see modified quasi-stationary method microwave oven, holes 291 MRI. see magnetic resonance imaging microwave power drying multistage drying processes 66 – moisture loss 293–295 multistage heat pump 349 – temperature distributions 295–299 MWCD. see microwave convective drying microwave power level, thermal runaway 298 MWFD. see microwave freeze-drying microwave spouted bed drying (MWSD) 280, MWSD. see microwave spouted bed drying 291 MWVD. see microwave vacuum-drying microwave-vacuum dryers 287–289 – industrial-scale 289 n microwave vacuum-drying (MWVD) 279 N2, feed foaming 169, 175 – rotary drum dryer 288 NaCl concentration, penetration depths 285 microwaves NaCl transport, liquid-solid systems – penetration depths 285 drying 255 366j Index nano XP emitters, electric radiatiators 333 o natural convection solar grain dryer 325 onions, quality in catalytic drying 331 nitrogen. see N2 opaque body 320 NO2, feed foaming 169 orange peel, freeze-drying 153 NOx emission, effect of tailpipe length 31 organic solvents and cosolvents, use in freeze- noise drying 144–146 – generated by pulse combustion 31 oscillating gas stream – intense with freeze-drying 265 – and atomization 37 – noise level, modeling of 42 43 – heat transfer coefficient in 38–42 non-thermal permeabilization of cell – modeling of 43–47 membranes, PEF technology 195 oscillations – nozzle arrangements, heat transfer in 10 11 – particles from ultrasonic waves 191 nozzle arrays – in PCD 27–28, 45 – characteristics 1 oscillatory electric charges 318 – compared to perforated plate 21 osmosis, ultrasound-assisted 216–217 – fl for drying at products 1 osmotic dehydration – drying principle 2 – combined drying processes 217–218 – – heat transfer in 4, 9, 11 12 – mass transfer barriers 205–206 – – in-line arrangement 8 11 oxygen-free drying environment 57, 74, 79 – staggered arrangement 8–11 – types of, 3 p nozzle diameter Page equation 343 – fl in uence on air temperature 20 paper drying – fl in uence on drying rate 22 – by nozzle arrays 1 nozzle distance, relative 6 – PCD 39–40 – and heat transfer 9 parameter optimization 227 nozzle-distributor spray-dryer 50 parenchyma cells, plant tissue composition 192 fi nozzle elds particle dynamic analysis (PDA), in PCD – – arrays of single nozzles 7 12 system 41 – – cylindrical bodies 14 16 particle phase, discrete (Lagrangian) – – design of 17 23 consideration of 107 – – hole channels 12 13 particle size influences, on contact – – perforated plates 13 14 heating 115–126 – fi – speci c energy consumption 18 19 particle velocity distributions, fluidized-bed fl nozzle-surface distance, in uence on simulations 109 perforated plates 23 particles NTU. see number of gasside transfer units – discrete particle modeling (DPM) 107–111 nucleation – foam-dried 166 – effect on tissue structures 227 – inert metallic 126–127 – freeze-drying 209 – moisture content of 47 –– – regulation of temperature 140 143 – single-particle drying curves 122 number of gasside transfer units (NTU) 103 – structure of foamed 165 numerical optimization techniques, PCD. see pulse combustion drying microwave drying 308 PDA. see particle dynamic analysis Nusselt functions peak wavelength, infrared radiation 321 – for perforated plates 13 PEF. see pulsed electric fields – – for single nozzles 16 17 penetration depth 285 Nusselt number – low, in infrared drying 329 – in array of individual nozzles 8 pepper tissue, osmotic dehydration 205 – for cylinders 16 perforated plates – local 9 – influence of pitch 23 – – PCD 40 41 – nozzle arrays 1, 3, 13–14 – of single nozzle, 6 –– specific energy consumption 21–23 Index j367 permeability, cell membranes 191 – micrographs 203 permittivity, relative complex 283 – PEF belt system 197 persimmon, drying kinetics modelling 260 – porosity 202 PFR. see plug flow reactor – viscoelastic properties 198 pharmaceutical industries, freeze-drying powders intensification 131–156 – caseinate 166 phase change, of water 237 – cottage cheese 172–173 phase controller, microwave drying 299 – density of 181 photons, electromagnetic radiation 321 – “looseness” of 176 physicochemical changes, food materials 305 – maltodextrin, microscopic images piezoelectric effect 244 166 piezoelectric transducers 243–244 – milk, foam spray-drying 167 pillow-shaped expansion, quality changes 306 – spray-dried (see spray-dried powders) pitch, influence for perforated plates 23 power drying Planck’s law 322 – moisture loss 293–295 plant raw materials – temperature distributions 295–299 – combined drying processes 216–217 – temperature variations 296–298 – freeze-drying 208–211 power ultrasound 237–278 – PEF technology 195–199 Prandtl number, PCD 40 – pretreatment 204–216 predehydration effects 348 – structural characterization of 192–204 pressure-gain valveless pulse – ultrasound 196–200 combustor 35 plant tissues pressure trace, in pulse combustor 29 – freezing tolerance improvement 211 pressurized chamber, pulse combustion dryer – PEF 195–199 with 51 – structure of 192 pretreatment – ultrasound 200 – air drying 207 plasma discharge problems, microwave- – plant raw materials 204–216 assisted drying 303 – quality characteristics 211–216 plasmalemma, plant tissue drying 192 – ultrasound 218 Ò plug flow reactor (PFR), ProCell units primary cyclone, of Pulse-Dryer ,52 103–106 primary freeze-drying stage 134 pneumatic dryers, pulse combustion 51 – chamber pressure in 136–137 polar molecules, microwave heating 282 – exergy losses during 136–138 polychromatic radiation 321 prismatic spouted beds 91–93 polyethylene baskets 289 – gas volume flow rate 97 polymer polyvinyl chloride (PVC), contact ProCell units heating 114 – bed pressure drop 98 poly(tetrafluoroethylene) (PTFE), contact – mass and heat transfer in 100–107 heating 114 – modeling of 108–111 pore formation, freeze drying 209 – Schmidt number 106 pore network, effects of soundwave 202 – Sherwood number 103–106 pores, effects on ultrasound drying 267 – single-particle correlations 107 pork loin brining, liquid-solid systems – stable operating range 99 drying 254–255 – symmetrical 95 porosity levels, plant tissue drying 192 process control, microwave-assisted porous matarials, as detergents 175–180 drying 308–310 porous matrix, freeze drying 208 process intensification. see also specific drying porous structure formation 350 processes/techniques potato chips, SSD 61–64 – process streams 250 potato slices, dielectric changes 304 – process variables 250–272 potato tissue – strategy for 309 – diffusivity 248 processing times 238 368j Index product characteristics, potential, of tailor – system components 195 made products 215 – technology 195–196 product quality pulsed electric fields (PEF) performance, – catalytic infrared drying 331 dependence on tissue structure 196 – microwave-assisted drying 279–316 pulsed electric fields (PEF) pretreatment – microwave drying 280 204–216 product temperature – effect on product quality 227 – as function of drying rate 18 – quality parameters 212–213, 214 – measurement of 286 pulsed electric fields (PEF) treatment, n-propyl acetate, retention of, foam spray- synergetic effects with ultrasound 228 drying 171 pulsed mode microwaves 298 PTFE. see poly(tetrafluoroethylene) purple light 302 pulse combustion PVC. see polymer polyvinyl chloride – aerodynamics 36–42 – versus continuous combustion 29–30 q – definition 27 quality changes, food 305–308 – heat and mass transfer 38–42 quality characteristics – noise generated by 31 – microwave drying 280 – principle of 27–32 – PEF pretreatment 211–216 pulse combustion drying (PCD) 27–53 – ultrasound combined drying 224–226 – driving force of 38 quality improvement, of foods and – liquid atomization in 37 biomaterials, by low-pressure SSD – modeling of 42–48 (LPSSD) 76–82 – temperature and velocity oscillations Quarter-Wave combustor 32 in 45 quartz lamps 332 – water evaporation in 41–42 pulse combustion flash dryers, for wood r waste 51 radiation drying pulse combustion pneumatic dryers 51 – classification, research, and applications pulse combustion pressurized-chamber of 323–332 dryers 51 – infrared 317 pulse combustion spray-dryers radiation heat transfer 318–323 – basic differences to classical spray-dryers 50 radiation intensity, slab thickness 328 – streamlines pattern in 44 radiation law 318 – typical configuration of 49 radiation properties, change during pulse combustors drying 340 – with aerodynamic valves 34–35 radiation spectrum 319 – configurations 48–53 radiative heat exchange 338 – design and operation 32–36 radiator types – with flapper valves 28, 33–34 – electric 333–335 – frequency-tunable 35–36, 48 – gas-heated infrared 335–337 – with mechanical valves 32–34 – infrared drying 332–337 – operation of 28, 35 rapid rate drying period, microwaves – pressure-gain valveless 35 293 – pressure trace in 29 rapid switching currents, ultrasonic – with rotary valve 32 equipment 243 – typical emissions from 30 raw materials, structural characterization Ò Pulse-Dryer , basic configuration of 52 192–204 pulsed electric fields (PEF) 191 Re-G-Ar-plot 99–100 – effects on freeze drying 209 reciprocity theorem 338 – freezing 208–211 red bell pepper, osmotic dehydration 205 – impact of 192–204 Reed-type valves, pulse combustors 33 – plant tissue structure 196–199 reference temperature, freeze-drying 134 Index j369 reflection shape factor, radiative heat exchange 338 – infrared drying 319–321 Sherwood number – of ultrasonic waves on interface 259 – PCD 40–41 reflectivity, infrared drying 338–339 – ProCell units 103–106 rehydration shockwaves, in tissue from ultrasound – contact ultrasound 226 waves 200 – quality parameters 212 short-wave burners, gas-fired infrared relative complex permittivity 283 source 336 relative nozzle distance 6 short-wave infrared emitter, electric resistance radiatiators 333 – increased, to water transport 293 shrimps, SSD 65–69 – kinetics modeling 257 shrinkage, food quality changes 305–306 – to vapor flow 140 simulations. see modeling Reynolds number single nozzles – for cylinders 16 – arrays of 3, 7–12, 21 – exponent of, 7 – heat transfer in 8 – PCD 40–41 – impinging jet drying 4–7 – spouted beds 94, 96, 99–100 – local heat transfer coefficients, 5 Rijke-type combustor 32 – Nusselt functions for 16–17 ring sonotrode 247 – Nusselt number of, 6 rotary kiln, with frequency-tunable pulse single-particle correlations, ProCell units 107 combustor 48 single-particle drying curve 122 rotary valves, pulse combustors 32 siren systems, ultrasonic equipment 243 – features 33 slab thickness, radiation intensity 328 rotational state, infrared energy absorption 319 slice thickness, effect on microwaves 294 SMD. see Sauter mean diameter s snack products 281 Salmonella anatum, inactivation under sodium caseinate, foamed 169–170 LPSSD 77–78 sodium chloride concentration, penetration salt. see also sodium chloride, NaCl depths 285 – addition during microwave drying 294 sodium chloride transport, liquid-solid systems – osmotic dehydration 205 drying 255 salt solution, evaporation of 46 solar drying 325–326 sample overheating, contact ultrasound 223 solar tunnel dryer 325 sample scanning, x-ray microtomography 328 solid food matrices, drying of 191 Saskatoon berries solid-liquid systems, mass transfer rate in – antioxidant properties of 150–151 drying 238 – optimum drying conditions 289 solids, contact ultrasound 216 saturated brine, acoustic pressure 258 solutes distribution, plant tissue 215 saturation pressure 2–3 solutes transport, liquid-solid systems sauce industry 243 drying 253 Sauter mean diameter (SMD) 37–38 solutions, containing sugar or salt 205 scattering 320 solvents, organic, use in freeze-drying Schmidt combustor. see Quarter-Wave combustor 144–146 Schmidt number sonication time reduction 221 – PCD 40–41 sonoluminescence, in tissue from ultrasound – ProCell units 106 waves 200 “second-order” phase transitions 122 sonotrodes 245 secondary atomization 38 – tip vibration 200 secondary effects. see specific energy sound pressure level (SPL), pulse consumption combustion 31 – ultrasound efficency measurements 252 soy protein isolate, additive for banana sensors, electroacoustic 252 foams 186 370j Index specific acoustic impedance 240 streamlines pattern, in pulse combustion specific drying energy 4 spray-dryer 44 specific energy consumption (SEC). see also stress-strain curves, tissue decomposition energy consumption 198 – as function of drying rate 18–22 stripping, of boundary layer 38 – microwave-assisted drying 300, 303 structural change stages, foods 306 specific moisture uptake capacity, of air 113 sublimation, frozen food materials 282 spectral distribution 321 sublimation flux spectrum – AFD 148 – infrared 321–323 – freeze-drying 133, 135 – radiation for drying 324 –– primary stage 138–139 spinach, freezing resistance 211 sublimation improvement, volumetric SPL. see sound pressure level heating 290 sponge effect 23, 242 sublimation interface 134 – contact ultrasound 218 – AFD 146, 148–149 – gas-solid system drying 256 sublimation rate, freeze-drying 139–140, – sample composition dependence 269 143–144 spouted bed drying, microwave 280, 291 – microwave-assisted 152 spouted beds sucrose – Archimedes number 94, 99–100 – addition during microwave drying 294 – bed pressure drop 98 – foamed 170–171 – different types of 86–93 – freeze-dried 144–145 – discrete particle modeling (DPM) 107–111 sugar, osmotic dehydration 205 – drying equipment 291–292 sugar beet, air drying 207–208 – flow patterns 88 sugar beet pulp, SSD 63–64 – gas velocity 86, 94 supercooled water 210 – gas volume flow rate 94–95, 97 superficial gas velocity, spouted beds – operating characteristics of 93–100 94–95 – Reynolds number 94, 96, 99–100 superheated steam-drying (SSD) 57–82 spray-dried powders – atmospheric-pressure 61–69 – microscopic images 166, 179–180 – drying rate 59 – non-foamed 172 – inversion phenomenon 59–60 spray-drying – low-pressure 69–76 – AFD 147 – principle of 58–61 – basic differences between classical and pulse surface moisture transfer 207 combustion 50 surface roughness 320 – foams 163, 169–170 surface temperature, gas burners 337 – typical configuration of 49 surfactants spray zone, of fluidized bed 107–111 – of biological character 182 SSD. see superheated steam-drying – foam stability 164 staggered arrangement, of nozzle arrays 8–11 synergetic effects, pulsed electric fields (PEF) stagnation point flows, definition 4 treatment and ultrasound 228 stagnation point region 5–7, 9, 17 staining methods, plant tissue drying 193 t starch granules, freeze drying 210 tailpipe 27 stationary fields 252 tailpipe length, effect on NOx emission 31 statistical significance, of experiments tank vibrations, ultrasonic bath 247 267–268 temperature steady-flow drying, difference to PCD 38 – of air 263–266 steam-drying, superheated 57–82 – effect on final product quality 182 Stefan-Boltzmann law 323, 338 – microwave heating 283 stepped-plate transducers 216 temperature change, microwave power – contact ultrasound 218 drying 299 Index j371 temperature control transducer types 243 – microwave drying 299 – with extensive radiators 248 – microwave power drying 298–299 – ultrasound drying 243–250 temperature distributions – ultrasound efficency measurements 251 – microwave-assisted drying 295–296 transient cavitation 241 – thin chips 296 transmembrane potential 196–197 temperature measurement systems 295 transmission, infrared drying 319–321 temperature modeling, carrots 298 transmission medium, ultrasound waves 202 temperature oscillations, in PCD 45 transmissivity, infrared drying 338–341 temperature overrun 298 transparent body 320 temperature runaway, danger of in microwave transport kinetics increase 254 drying 281 tubes, immersed, for contact heating 112, 125 temperature sensors, fiber-optic 286 tunnel dryer, freeze-drying 149 temperature stages, constant microwave turbulence level, process intensification 250 power 296–297 turbulences, ultrasonic field reduction 259 temperature variations turgor, plant tissue drying 192 – microwave power drying 296–298 twin-tube emitters, electric radiatiators – variable microwave power drying 333–334 298–299 tert-butyl alcohol, as solvent in freeze- u drying 144–146 ultrasonic bath 247 textiles, drying by nozzle arrays 1 – acoustic pressure variations 260 texture, impact of PEF on 197–198 ultrasonic energy thermal effects, water evaporation 224 – deployed in sample 251 thermal radiation 318 – source 243 – interaction with material 320 ultrasonic equipment 242–250 thermal runaway 298 – for gaseous media 249 thermal treatment 194 – for liquid applications 246 thermocouples ultrasonic field – metal-shielded 295 – chaotic 247 – temperature measurements 330 – measurement of 251 thin-layer solar drying 325 ultrasonic field effects, parameters 251 threshold ultrasonic intensity 239, 252 – for effects of ultrasound 221 – effects during drying 253 – ultrasonic intensity 267 – medium characteristics influence 258 tiles, drying by nozzle arrays 1 – threshold 267 time-adjusted power 299 – zones 252 tip vibration, ultrasound waves 201 ultrasonic power, drying applications tissue integrity, freeze drying 208 250–263 tissue structure ultrasonic radiation effects, temperature – analysis of 192–195 dependence 263 – pulsed electric fields (PEF) 195–199 ultrasonic system components 219 tissues ultrasound 199–204 – chemical changes 212 – air drying 218–221 – contact ultrasound effects 204 – combined drying processes 216–217 – drying by nozzle arrays 1 – effects in liquids 225 tomato paste, foam-mat drying 184 – effects on mass transfer 241–242 tonoplasts, plant tissue drying 192 – freeze-drying 221–224 tortilla chips, SSD 61–62 – high-intensity 199 total emissive power 321 – impact of 192–204 total emissivity, different surfaces 323 – intensity of 252 total energy, ultrasonic system 239 ––mass transfer effects 241 transducer design 222 – intensity threshold 267 372j Index – low-intensity 199 velocity oscillations, in PCD 45 – osmotic dehydration 217–218 VFD. see vacuum freeze-drying – plant tissue structure 200–204 vibrating drying chambers 216 – power ultrasound 237–278 vibration characteristics, ultrasound – process variables 250–272 drying 219 – quality characteristics 224–226 vibrational state, infrared energy – synergetic effects with PEF 228 absorption 319 ultrasound-assisted drying viscous flow, ultrasound effects 241 – freeze-drying 152–153 vitamin C. see ascorbic acid – power ultrasound 237–278 volatile aroma compounds preservation ultrasound-assisted osmosis 217 307 ultrasound drying volume changes, food materials 305, 307 – effects of 252 volume flow rate, of heated air 1 – effects of sample composition 266–268 volumetric heating, microwave drying 280 ultrasound pretreatment 218 ultrasound-sample interaction 266–272 w ultrasound treatment, intermittent 220, 225 wall-to-bed heat transfer 118 ultrasound waves 239–241 water, effective diffusivity 253 – power 239–240 water distribution, quality parameters 215 – sonotrode tip vibration 201 water evaporation – water removal improvement 202–203 – in PCD 41–42, 46 uniform expansion, quality changes 306 – thermal effects 224 water films, spectral absorption 340 v water removal vacuoles, plant tissue drying 192 – improvement 202 vacuum, effect on microwaves 294 – under vacuum 221 vacuum chamber, integration of ultrasound water transport, liquid-solid systems transducer 223 drying 253 vacuum-drying (VD) wave damping, ultrasound 191, 199 – compared to LPSSD 70–76 wave frequency, ultrasound parameters – equipment 287–290 251 – FIR-assisted (VACUUM-FIR) 76–82 wavelength, ultrasonic system 240 vacuum freeze-drying (VFD) 237 wet formulation, fluidized-bed processes 85 – of liquids 139–146 whey protein concentrate, additive for banana – ultrasound assisted 265 foams 186 vacuum impregnation 211 whistles, ultrasonic equipment 243 vacuum rotary drum dryer 288 Wien’s displacement law 322 vacuum seal 288 wood veneer, drying by nozzle arrays 1 valves. see specific types wood waste, dried in pulse combustion flash vapor flow, resistance to 140 dryer 51 vapor pressure, at interface of working width, electric radiatiators 333 sublimation 140 Wurster equipment, spouted bed 90 vapor transfer, in PCD 42 – modeling of 108–111 vaporization, enthalpy of 3–4 vavuum drying (HAD), bacteria x inactivation 77 X-ray microtomography, nondestructive VD. see vacuum-drying sample scanning 328 vegetable drying (VD) 192 – cryoprotectants usage 211 y – microwave-assisted 150–151 “Yankee dryers” 47