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School of Chemical Engineering and Industrial Chemistry The University of New South Wales

Catalytic Distillation for the Synthesis of Tertiary Butyl

by

Tomasz Safinski

A thesis submitted in partial fulfilment of the requirements for the degree of

Doctor of Philosophy

2005

Certificate of Originality

I hereby declare that this submission is my own work and to the best of my knowledge it contains no material previously published or written by another person, nor material which to a substantial extent has been accepted for the award of any other degree or diploma at UNSW or any other educational institution, except where due acknowledgment is made in the thesis. Any contribution made to the research by others, with whom I have worked at UNSW or elsewhere, is explicitly acknowledged in this thesis.

I also declare that the intellectual content of this thesis is the product of my own work, except to the extent that assistance from others in the project’s design and conception or in style, presentation and linguistic expression is acknowledged.

(Signed)......

ii

Dedicated to

Aoife

iii Abstract

Catalytic Distillation for the synthesis of tertiary butyl alcohol (TBA) is investigated in this thesis. The solvent, glycol, is proposed as a means of overcoming challenges, which limit the potential benefits of the application of reactive separation for TBA. The proposed action of the solvent is that of extractive distillation entrainer, thus a new unit operation of Catalytic Extractive Distillation (CED) is suggested.

The solubility of in water, TBA, and ethylene glycol and their binary and ternary mixtures, at different temperatures, is measured and correlated. The solubility is found to be highly non-linear in solutions containing TBA. The kinetics of isobutylene hydration over Amberlyst 15 is characterised in the presence of ethylene glycol. The solvent is found to promote reaction rate, however it is also found to compete for reaction with isobutylene. Water is found to strongly inhibit the reaction of ethylene glycol and isobutylene. The selectivity ratio of TBA to by products is determined and found to improve with increased temperature and lower solvent concentration.

Bale packing is chosen as catalytic distillation hardware for the containment of Amberlyst 15 and its two-phase fluid dynamics characterised for the first time. Raschig rings are used as a benchmark for the study. Bale packing is found to exhibit two ranges of backmixing behaviour in the pre-loading regime. This behaviour is attributed to the three levels of porosity of the hardware and indicative of low rates of catalyst/liquid renewal.

The effectiveness of ethylene glycol as extractive distillation entrainer for the separation of the TBA/water azeotrope over Bale packing is investigated and the solvent found to be highly effective.

The mass transfer resistances to isobutylene transport are determined for countercurrent fixed bed reactor (CFBR) application of Bale packing. It is found that ethylene glycol improves mass transfer coefficients attainable. Catalytic Extractive Distillation is

iv implemented over Bale packing and the ability of the solvent to improve reaction rates and purity of TBA demonstrated. However, the reaction rates achieved have much scope for improvement through increased isobutylene availability.

In response to poor liquid renewal of static packing such as Bale packing and the necessity of improved isobutylene transport a new form of catalytic distillation reactor design is proposed, the Basket Impeller Column (BIC). The BIC combines the mass transfer benefits of a rotating basket reactor with that of a dual flow column. Capacity of the new hardware is determined and correlated. Separation and reactive separation are demonstrated to be feasible. It is found that Damköhler number can be varied directly using the additional process variable of speed of rotation.

v List of Publications

T. Safinski, A. A. Adesina, Axial Mixing in Two-Phase Flow Countercurrent Operation of a Trickle Bed Reactor, 6th International Conference on -Liquid and Gas-Liquid- Solid Reactor Engineering, Vancouver, BC Canada, August 17-20 August, 2003

T. Safinski, A. A. Adesina, B. R. Young, Parametric Study of a Novel Basket Impeller Column, 6th International Conference on Gas-Liquid and Gas-Liquid-Solid Reactor Engineering, Vancouver, BC Canada, August 17-20 August, 2003

T. Safinski, A. A. Adesina, B. R. Young, Parametric Study of a Novel Basket Impeller Column, The Canadian Journal of Chemical Engineering, 81, 574-580, 2003

T. Safinski, A. A. Adesina, Development of a Novel Basket Impeller Dualflow Tray Catalytic Distillation Reactor, Industrial Engineering and Chemistry Research, Ind. Eng. Chem. Res. 44, 6212-6221, 2005

vi Acknowledgments

I would like to thank my supervisor, Prof Adesoji A Adesina, for his advice, guidance, and patience throughout the period of this research. I am greatly indebted to him for his selfless help and effort. I would like to thank him for this wonderful research opportunity.

I would also like to thank:

Dr Chiming Zhang, for providing technical advice and practical instruction.

Steve Preece, Scientific Illustrations UNSW, for his technical advice regarding analysis of digital images.

Paul Brockbank and Phil Thompson, CEIC Workshop, for all their high quality work.

John Starling, CEIC Professional Officer, for chasing orders and persistence on GLP.

All the Honours, Masters, PhD, and Industrial Training students; and visiting academics who have shared ideas, experience and companionship, while working as part of our research group. Particular thanks, to students who worked on projects in my care, including: Huy Tuan Nguyen, Catherine Vagi, Sonny, Edwin Mandagi, Eko, Samuel Sze, Daniel Farmansyah, Robert and Han Cheng Hsu.

Faculty of Engineering UNSW, for their financial support in granting me a Faculty Scholarship.

I would like to thank my wife, Aoife, for her constant love, support and encouragement.

vii Table of Contents

Title Page...... i Certificate of Originality...... ii Abstract...... iv List of Publications...... vi Acknowledgments...... vii List of Figures...... xiv List of Tables...... xvii

1 Introduction...... 1 1.1 Literature Cited...... 4 2. Literature Review...... 6 2.1 TBA For The Purpose Of Blend Additive...... 6 2.1-1 Octane Number...... 6 2.1-2 Fuel Additives, Oxygenates and Others, History of Developments...... 8 2.1-3 Fuel Oxygenates, the Global Situation...... 9 2.1-4 The Australian Situation...... 10 2.1-5 TBA as Fuel Oxygenates...... 10 2.2 Industrial Production of TBA...... 13 2.2-1 TBA as a By-product of the Halcon Propylene Oxide Process...... 13 2.2-2 Direct Hydration of Isobutylene...... 13 2.3. Thermodynamics and Kinetics of the Hydration of Isobutylene...... 15 2.3-1 Thermodynamics of Isobutylene Hydration...... 15 2.3-2 Homogeneous Kinetics of Isobutylene Hydration...... 16 2.3-3 Heterogeneous Kinetics of Isobutylene Hydration...... 18 2.3-4 Heterogeneous Kinetics of Similar Systems...... 25 2.3-5 Use of Solvents for Isobutylene Hydration...... 29 2.4 Extractive Distillation Technology...... 31 2.4-1 Entrainer Selection...... 33 2.4-2 The TBA and Water Azeotrope...... 34 2.4-3 Extractive Distillation Applications...... 35 2.5 Catalytic Distillation Technology...... 37

viii 2.5-1 New Applications and Process Improvement ...... 38 2.5-1-1 Methyl tert-Butyl Ether, MTBE ...... 40 2.5-1-2 Ethyl tert-Butyl Ether, ETBE...... 48 2.5-1-3 tert-Amyl Methyl Ether, TAME...... 49 2.5-1-4 tert-, TAA...... 50 2.5-1-5 Dehydration of TBA...... 51 2.5-2 Process Synthesis and Design Tools...... 52 2.5-2-1 Catalytic Distillation Thermodynamics...... 53 2.5-2-2 Limiting Regimes, Damköhler and Hatta Numbers...... 57 2.5-3 Modelling and Simulation...... 59 2.5-3-1 Equilibrium Based Models...... 59 2.5-3-2 Non-Equilibrium Based Models...... 61 2.5-4 Catalytic Distillation Hardware Design...... 62 2.5-4-1 Catalyst Containment...... 64 2.5-4-2 Catalyst Alteration...... 66 2.5-5 Dynamics and Control...... 67 2.6 Catalytic Extractive Distillation...... 68 2.7 Countercurrent Fixed Bed Reactors...... 71 2.7-1 CFBR Applications and Technology...... 71 2.7-1-1 Comparing CFBR to TBR...... 73 2.7-2 Fluid Dynamics...... 74 2.7-2-1 Liquid Phase Mixing...... 74 2.7-2-2 Gas Phase Mixing...... 76 2.7-3 CFBR Analysis ...... 76 2.7-3-1 Basic Reactor Treatment...... 77 2.7-3-2 Integral Reactor Treatment...... 77 2.7-3-3 CFBR Mass Transfer Correlations...... 77 2.8 Literature Cited...... 79 3. Proposed Process and Investigation...... 90 3.1 Towards Catalytic Distillation...... 90 3.2 Proposed Process, Solvent Selection ...... 95 3.3 Selection of Catalytic Distillation Reactor ...... 102 3.4 Proposed Investigation ...... 103

ix 3.5 Literature Cited ...... 105 4. Kinetics of the Hydration of Isobutylene to TBA over Amberlyst-15 in the Presence of the Solvent Ethylene Glycol ...... 109 4.1 Background Theory...... 111 4.1-1 Mass Transport...... 111 4.1-2 Kinetic Models...... 115 4.2 Experimental...... 119 4.2-1 Stirred Basket Reactor...... 119 4.2-2 Catalyst Pretreatment...... 123 4.2-3 Analytical Procedure ...... 124 4.2-4 Experimental Procedure...... 126 4.2-5 Choice of Variables and Experimental Design...... 132 4.2-6 Calculation Procedure...... 133 4.3 Results and Discussion...... 134 4.3-1 Isobutylene Solubility...... 134 4.3-2 Transport Analysis...... 144 4.3-3 Pure Water Runs...... 148 4.3-4 Empirical Model for Ethylene Glycol Mediated Isobutylene Hydration. 150 4.4 Heterogeneous Kinetic Models...... 158 4.4-1 Proposed Models...... 158 4.4-2 Model Evaluation...... 160 4.5 Concluding Remarks ...... 168 4.6 Nomenclature ...... 170 4.7 Literature Cited...... 172 5. Bale Reactor Characterisation...... 175 5.1 Background...... 176 5.1-1 Packed Column Capacity...... 176 5.1-2 Reactor Fluid dynamics and Residence Time Distribution...... 182 5.2 Bale and Raschig Ring Packing...... 186 5.2-1 Bale Packing Preparation Procedure...... 186 5.2-2 Packing Characterisation...... 188 5.3 Capacity of Bale Packing...... 189 5.3-1 Experimental...... 189

x 5.3-2 Results and Discussion...... 193 5.3-2-1 Pressure Drop...... 193 5.3-2-2 Liquid Holdup...... 194 5.4 Gas Phase Mixing, CFBR Flow Conditions: Pre-Loading Regime...... 198 5.4-1 Experimental, Gas Phase...... 199 5.4-1-1 Automation and DAQ with LabVIEW...... 200 5.4-1-2 CFBR Gas Side Fluid Dynamics Operating Procedure...... 202 5.4-1-3 Experimental Design ...... 203 5.4-1-4 Analysis and Calculation Procedure...... 205 5.4-2 Gas Phase Fluid Dynamics Results and Discussion...... 210 5.5 Liquid Phase Mixing in the Pre-Loading Regime...... 227 5.5-1 Experimental...... 227 5.5-1-1 Liquid Phase Procedure...... 228 5.5-1-2 Experimental Design...... 231 5.5-1-3 Calculation Procedure...... 231 5.5-2 Liquid Phase Fluid Dynamics Results and Discussion...... 234 5.6 Concluding Remarks ...... 246 5.7 Nomenclature ...... 249 5.8 Literature Cited...... 251 6. Extractive Distillation over Bale Packing...... 253 6.1 Governing VLE ...... 255 6.1-1 Wilson Model and Parameters...... 255 6.1-2 UNIFAC Model and Parameters...... 257 6.1-3 VLE Predicted...... 256 6.1-4 Relative Volatilities and Selectivity...... 264 6.1-5 Minimum Solvent Required...... 267 6.2 Experimental...... 270 6.2-1 Distillation Equipment...... 270 6.2-2 Analytical Procedure ...... 272 6.2-3 Experimental Procedure...... 273 6.2-4 Experimental Variables...... 278 6.3 Extractive Distillation Results...... 279 6.3-1 Range of Operation...... 282

xi 6.3-2 Temperature Profiles ...... 282 6.3-3 Separation by Extractive Distillation...... 284 6.3-4 Theoretical Plates Achieved...... 287 6.4 Towards CED, Further Considerations...... 295 6.4-1 The Effect of Isobutylene Upon Selectivity...... 295 6.4-2 Matching Separation and Reaction Rates over Bale Packing...... 298 6.5 Concluding Remarks...... 300 6.6 Nomenclature ...... 301 6.7 Literature Cited...... 303 7. Synthesis of TBA under CFBR and CED Conditions in the Presence of Ethylene Glycol ...... 305 7.1 Three Phase Bale Reactor Mass Transfer...... 306 7.2 Synthesis of TBA over Bale Packing, CFBR Conditions...... 309 7.2-1 CFBR Experimental...... 309 7.2-1-1 Analytical Procedure ...... 309 7.2-1-2 CFBR Procedure...... 310 7.2-1-3 Experimental Design ...... 314 7.2-1-4 Calculation Procedure...... 316 7.2-2 Results of Single Pass CFBR Runs...... 319 7.2-3 Discussion of Single Pass Study...... 322 7.2-4 Results of CFBR Recirculation Runs...... 330 7.2-5 Analysis and Discussion of Recirculation Study...... 332 7.3 Synthesis of TBA by CED over Bale Packing...... 341 7.3-1 Experimental...... 342 7.3-1-1 CED Procedure...... 342 7.3-1-2 Experimental Design...... 347 7.3-2 Results and Discussion...... 348 7.4 Concluding Remarks ...... 358 7.5 Nomenclature ...... 358 7.6 Literature Cited...... 360 8. Basket Impeller Column...... 362 8.1 Literature Relevant to BIC Development...... 364 8.1-1 Dual-Flow Columns...... 365

xii 8.1-2 Weeping and Bubble Formation at a Submerged Orifice...... 369 8.1-3 Rotation and Distillation...... 370 8.2 BIC Single Stage Evaluation...... 372 8.2-1 Single Stage Apparatus ...... 372 8.2-2 Single Stage Procedure...... 374 8.2-3 Experimental Design ...... 377 8.2-4 Single Stage Sieve Tray Study, Results and Discussion...... 379 8.2-5 Single Stage Fluid Dynamics Study Results and Discussion...... 382 8.2-5-1 Liquid Phase Holdup ...... 382 8.2-5-2 Gas Phase Holdup...... 386 8.2-5-3 Froth Height...... 390 8.2-5-4 Pressure Drop...... 391 8.2-5-5 Power Requirement...... 393 8.3 BIC Multistage Case Studies...... 396 8.3-1 Multistage Column Apparatus...... 396 8.3-2 Case Study A. System Characterisation...... 399 8.3-2-2 Fluid Dynamics under Separative Conditions...... 399 8.3-2-3 Overall Column Efficiency...... 405 8.3-3 Case Study B. Application to Catalytic Distillation...... 408 8.3-3-1 TBA Dehydration Kinetics...... 408 8.3-3-2 Catalytic Distillation Runs...... 414 8.4 Concluding Remarks and Further Considerations...... 419 8.4-1 Conclusion...... 419 8.4-2 Potential Applications of the BIC...... 419 8.5 Nomenclature ...... 421 8.6 Literature Cited...... 423 9. Conclusion and Recommendations...... 427 Appendix A1: Materials and Specifications...... 432 Appendix A2: Solubility and Kinetics Data...... 433 Appendix A3: Digital Image Processing...... 436 Appendix A4: RTD Data Smoothing and Processing...... 437 Appendix A5: UNIFAC Based VLE Calculation...... 440 Appendix A6: Wilson Model Based VLE Calculation...... 443

xiii List of Figures

Figure 2.01 Residue curves for the extractive distillation of and using either water or , Bauer and Stichlmair (1995)...... 34 Figure 2.02 Reactive residue curve for heterogeneously catalysed MTBE synthesis at operating pressure P = 0.8 MPa; Damköhler number (a) Da = 0 (b) Da = 10-04 (c) Da = 2 × 10-04 and (d) Da = 1, Thiel et al (1997)...... 47 Figure 2.03 Construction for graphical determination of the the loci of possible reactive azeotropes for a reaction system of the following form a+b = c ,

where π is the pole point defined by xiπ=νi/Σνj. Frey and Stichlmair (1999)...... 54 Figure 2.04 Equilibrium stage representation in most versatile form...... 60 Figure 2.05 NEQ stage representation showing mass and energy fluxes and the use of film theory...... 61 Figure 3.01 Equilibrium lines at T1 60oC, T2 80oC, T3 100oC and T4 120 oC...... 95 Figure 3.02 Determination of the likelihood of the occurrence of reactive azeotropes within the TBA system...... 94 Figure 3.03 CED Design A Reversal of Natural Separation Tendency of TBA and Water. Requiring a branched high molecular weight solvent such as 2,2,4-trimethyl-1,3-pentanediol of lower polarity than TBA...... 99 Figure 3.04 CED Design B Promotion of Natural Separation Tendency of TBA and water. Requiring a low molecular weight solvent such as ethylene glycol of polarity intermediate to that of water and TBA...... 100 Figure 4.01 Schematic of (a) the stirred basket slurry reactor employed, showing position of basket assembly and baffles (b) baffles and(c) baskets which the hold catalyst...... 122 Figure 4.02 Stirred Basket Reactor Rig for Kinetics Determination...... 129 Figure 4.03 The solubility of isobutylene, xIB, versus partial pressure yIBP and temperature for (a) water, (b) ethylene glycol and (c)TBA...... 136 Figure 4.04 The effect of temperature on the Henry’s constant obtained, showing an Arrhenius type behaviour...... 138 Figure 4.05 Determination Δho and Δso of isobutylene absorption in W/EG/TBA. Also plotted are the data of Leung et al (1987) for comparison...... 138 Figure 4.06 Binary mixtures of water, EG and TBA: (a)W/EG, (b) W/TBA and (c) EG/TBA, demonstrating the strong positive enhancement effect of TBA upon isobutylene solubility expressed here in mol.L-1...... 143 Figure 4.07 The elimination of the effect of external mass transport artefacts upon -1 -1 observed reaction rate, rTBA (mol.g .s ), through adequate stirring speed, Ω (rpm).... 144 Figure 4.08 The determination of internal mass transfer parameters for wet Amberlyst 15

of dp: 350, 450 and 550 μm at 361 K...... 145 Figure 4.09 The reaction rate of TBA formation with isobutylene liquid phase concentration and temperature...... 148 Figure 4.10 Comparison of pseudo first order rate constants obtained in the current study to published data at temperatures particular to each...... 149 Figure 4.11 The influence of ethylene glycol upon isobutylene hydration at (a) xEG 0.32 (b) xEG 0.55 and (c) xEG 0.81...... 154 Figure 4.12 Formation rate of MET versus CIB for xEG = 0.81...... 156 Figure 4.13 Rates of TBA and MET formation (a) water/isobutylene concentration product in the presence of ethylene glycol (b) ethylene glycol/isobutylene concentration product in the presence of water...... 157 Figure 4.14 Heterogenous Kinetic Model Further Evaluation of

xiv (a) M5 and (b) M7 parity plots...... 167 Figure 5.1 Laboratory Scale Bale packing prepared...... 187 Figure 5.02 Detail of Reactive Zone Jacketed Reactor and distributors...... 191 Figure 5.03 Detail of (a) top Liquid Distributor and Gas Collector fabricated in Pyrex and; (b) bottom Gas Distributor and Liquid Collector fabricated in brass and copper...... 192 Figure 5.04 Pressure Drop of Dry and Wetted Bale Packing...... 193 Figure 5.05 Liquid Phase Holdup of Bale Packing...... 195 Figure 5.06 Loading and Flooding Points for Countercurrent operation of Bale Packing.... 196 Figure 5.07 Front Panel of the Program Blue.vi used for Automation and Data Acquisition for Gas Phase RTD Experiments...... 200 Figure 5.08 Rig for Gas Side Fluid Dynamics Study...... 201 Figure 5.09 Measure of Gas Phase Tail Suppression α as a Function of the Laplace Parameter T for (a) Raschig and (b) Bale packing...... 209 Figure 5.10 Exit Age Distributions, E(t/tm), for various ReG at ReL = 490 of (a) raschig and (b) Bale packing...... 213 Figure 5.11 Exit Age Distributions, E(t/tm), for various ReL at ReG = 56 of (a) raschig and (b) Bale packing...... 214 Figure 5.12 Mean residence time as a function of ReG at different ReL for (a) Raschig Ring and (b) Bale packing...... 216 Figure 5.13 Gas Phase Total Holdup as a function of ReG at different ReL for (a) Raschig Ring and (b) Bale packing...... 218 Figure 5.14 Gas Phase Reactor Peclet number, Per,G, as function of ReG for different ReL for (a) Raschig Ring and (b) Bale packing...... 220 Figure 5.15 Gas Phase Reactor Peclet number, Per,G, replotted as a function of ReL for different ReG for (a) Raschig Ring and (b) Bale packing...... 221 Figure 5.16 Parity plots of Equations (5.55) and (5.56) modelling Gas Phase Reactor Peclet number, Per,G, for (a) Raschig Ring and (b) Bale packing...... 224 Figure 5.17 Rig for Experimental Determination of CFBR Liquid Side Fluid Dynamics... 229 Figure 5.18 Example Images and Values obtained...... 232 Figure 5.19 Calculation example corresponding to images of Figure 5.18, Bale packing with ReL = 735...... 233 Figure 5.20 Liquid Phase lack of dependence of Per,L upon ReG for (a) Raschig Ring and (b) Bale packing...... 236 Figure 5.21 Liquid Phase Exit Age Distributions as a Function of ReL for (a) Raschig Ring and (b) Bale packing...... 237 Figure 5.22 Liquid Phase tm as a function of ReL for (a) Raschig Ring and (b) Bale packing...... 239 Figure 5.23 Liquid Phase Dynamic Holdup, HD,L as a function of UL for (a) Raschig Ring and (b) Bale packing...... 240 Figure 5.24 Liquid Phase Reactor Peclet number, Per,L, variation in ReL for (a) Raschig Ring and (b) Bale packing...... 242 Figure 5.25 Raschig Ring Fluid or Packing Peclet number as a function of Reu,L and comparison with: Furzer (1972), Otake (1958), Bennett (1970), Sater (1966), Stockar (1984), and Macias-Salinas (2000)...... 244 Figure 6.01 Comparison of binary x-y data for the TBA water system generated using the Wilson and UNIFAC models showing the presence of an azeotrope occurring at approximately xTBA 0.60 to 0.70...... 262 Figure 6.02 X’-Y’, VLE data presented on a solvent free basis demonstrating the ability of ethylene glycol to cross the TBA/water azeotrope. Generated using (a) Wilson model with parameters of Liu et al. and (b) UNIFAC model...... 263 Figure 6.03 Relative volatilities predicted by (a) Wilson and (b) UNIFAC models...... 265 Figure 6.04 Selectivity calculated by the Wilson Model at x’TBA = 0.6754 and

xv UNIFAC at x’TBA = 0.6107, the respective predicted solvent free azeotropes...... 266 Figure 6.05 The approach of the equilibrium line to the 45o line for xr, (a) Wilson and (b) UNIFAC models...... 269 Figure 6.06 QuickFit Distillation Glassware: 1. Condenser Connection; 2. Reflux Section; 3. Solvent Feed Section; 4. Jacketed Bale packed Section; 5. Reboiler bypass; and 6. Reboiler...... 271 Figure 6.07 Extractive Distillation Rig...... 275 Figure 6.08 Extractive Distillation Stream Diagram...... 279 Figure 6.09 Extractive Distillation Temperature Profile at Az 0.57 mol.min-1...... 283 Figure 6.10 Purity gained by extractive distillation as a function of xEG for different azeotrope loading Az: 1. 0.49, 2. 0.57, 3. 0.71, and 4. 0.82 mol.min-1...... 285 Figure 6.11 Recovery in extractive distillation as a function of xEG for different azeotrope loading Az: 1. 0.49, 2. 0.57, and 3. 0.71 mol.min-1...... 285 Figure 6.12 Purity against Recovery for the extractive distillation study conducted with different azeotrope loading Az: 1. 0.49, 2. 0.57, and 3. 0.71 mol.min-1...... 286 Figure 6.13 Analysis of runs to for Az 0.49 mol.min-1 for experiment 1. xEG = 0.46, 2. xEG =0.52 and 3. xEG = 0.62...... 291 Figure 6.14 Analysis of runs to for Az 0.57 mol.min-1 for experiment 6. xEG = 0.42, 7. xEG = 0.54, 8. xEG = 0.63, 9. xEG = 0.66 and 10. xEG = 0.72...... 292 Figure 6.15 Analysis of runs to for Az 0.71 mol.min-1 for experiment 11. xEG = 0.37, 12. xEG = 0.48, 13. xEG = 0.61, 14. xEG = 0.68 and 15. xEG = 0.73...... 293 Figure 6.16 Extractive distillation Number of Theoretical stages achieved (Ntheoretical) with variation in xr and different azeotrope feed loadings...... 294 Figure 6.17 Effect of isobutylene upon the extractive distillation of water and TBA with ethylene glycol (a) x’-y’ curves in TBA and (b) selectivity as a function of xIB... 297 Figure 7.01 CFBR rig for isobutylene hydration over Bale Packing containing Amberlyst 15...... 311 Figure 7.02 CFBR liquid outlet concentration of product TBA for Single Pass Water Runs (a) ReG = 22 and (b) ReG = 11...... 320 Figure 7.03 CFBR liquid outlet concentration of product TBA for Single Pass runs using 30 mol % ethylene glycol...... 321 Figure 7.04 Observed TBA Reaction for (a) solvent absent water runs and (b) solvent present, ethylene glycol runs...... 322 Figure 7.05 Overall reaction rate observed and predicted...... 327 Figure 7.06 Mass Transfer Resistances to Isobutylene hydration over Amberlyst 15 as a function of ReL...... 328 Figure 7.07 Mass Transfer Resistances to Isobutylene hydration over Amberlyst 15 as a function of Per,L,B...... 329 Figure 7.08 A comparison of selectivity factors obtained in the Kinetic Study and CFBR runs, showing the run id of the experimental design...... 338 Figure 7.09 MET observed overall resistance as a function of observed TBA reaction rate, rTBA...... 339 Figure 7.10 CED Experimental Rig...... 344 Figure 7.11 Column Stream Number Specification...... 348 Figure 7.12 The effect of ethylene glycol feed rate upon observed TBA reaction rate under CED conditions and zero reflux...... 356 Figure 7.13 The effect of ethylene glycol feed rate upon observed TBA concentration under CED conditions and total reflux...... 356 Figure 8.01 Basket Impeller Column: (a) Holdup Determination and Feasibility Rig, (b) liquid distributor design & (c) basket impeller design...... 375 Figure 8.02 Normality plot for the experimental design of the first study, indicating significant factors and interactions...... 381 Figure 8.03 Liquid Holdup with Basket size and RPM at ReG 960, and ReL 12...... 382 Figure 8.04 The effect of RPM. Left to Right RPM 0, 50, 75, 100, 125 and 150 at

xvi ReG 960 and ReL 12...... 383 Figure 8.05 HL with ReG at RPM 100 and ReL 12...... 384 Figure 8.06 The effect of ReG. Left ot right ReG 265, 613, 960, 1309 and 1656 at RPM 100 and ReL 12...... 384 Figure 8.07 Parity plot for Equation (8.09)...... 385 Figure 8.08 HG with (a) RPM and basket size and (b) ReG and ReL...... 387 Figure 8.09 The ratio HGFroth/HL with RPM showing the greater extent of aeration and recirculation of bubbles...... 388 Figure 8.10 Parity plot for Equation (8.14)...... 389 Figure 8.11 Parity plot for Equation (8.15)...... 390 Figure 8.12 Pressure drop with (a) RPM and (b) ReG...... 391 Figure 8.13 Parity plot for Equation (8.16)...... 392 Figure 8.14 Power P (W) with RPM at ReG 1354 and ReL 23...... 393 Figure 8.15 Multistage Basket Impeller Column Setup...... 397 Figure 8.16 Multistage Basket Impeller Column Internals (a) Column Ends (b) Basket Cluster (c) Tray Support and Arrangement of One Stage (d) Tray Open Area...... 397 Figure 8.17 Froth height, hf, variation with basket speed for different trays, (a) 5.8 %OA and (b) 11.5 %OA...... 400 Figure 8.18 Parity plot between observed and predicted hf values...... 403 Figure 8.19 Influence of basket impeller speed on column pressure drop , ΔP (a) 5.8 %OA and (b) 11.5 %OA...... 404 Figure 8.20 Overall Column Efficiency, EOC, as a function of stirrer speed: (a) 5.8 %OA and (b) 11.5 %OA...... 406 Figure 8.21 Effect of stirring speed on TBA dehydration rate over Amberlyst-15...... 409 Figure 8.22 Kinetics of TBA dehydration rate...... 409 Figure 8.23 Concentration profile obtained at Ω=200 and 500 rpm for TBA feed rate of 10-03 mol s-1...... 416 Figure 8.24 Enhancement factor at various TBA feed rates...... 418 Figure 8.25 Damköhler number as a function of TBA feed rates...... 418 List of Tables

Table 2.01 Maximum concentration of compounds under EU and US legislation...... 10 Table 2.02 Physical Properties of Fuel Oxygenates, Drogos and Dias (2001)...... 11 Table 2.03 Composition of a typicalC4 Stream...... 13 Table 2.04 Solubility of isobutylene within mixtures of water and TBA...... 16 Table 2.05 Properties of Amberlyst 15 dry, Rohm and Haas (R&H)...... 18

Table 2.06 Internal mass transfer parameters: De, φ and η, for isobutylene hydration over Amberlyst 15, Leung et al (1986)...... 20 Table 2.07 Activation Energies reported for heterogeneous catalysis over ion exchange resins...... 22 Table 2.08 The kinetics of ethylene glycol etherification with isobutylene, at 60 oC, 1.5 atm, 10 wt % loading Amberlyst 15...... 31 Table 2.09 Insensitivity of water/TBA azeotrope towards pressure...... 34 Table 2.10 The selectivity of ethylene glycol for EtOH and TBA...... 35 Table 2.11 Potential Entrainers for the separation of water and TBA, Berg and Yang (1992)...... 35 Table 2.12 Predicted column dimension given a12, Berg and Yang (1992)...... 36 Table 2.13 Approximate values of HETP over Bale packing (1997)...... 65

xvii Table 2.14 CED for methyl acetate synthesis column properties...... 70 Table 2.15 Dispersion Model, Pe, Correlations for countercurrent operation over Raschig Rings...... 75 Table 2.16 Gas Phase Mixing in the presence of Liquid Flow, for Raschig Rings (RR), in reactor based parameters...... 76 Table 3.01 Relative volatility of water and TBA at infinitedilution for different solvents, of the glycol sereis, determined using UNIFAC...... 97 Table 4.01 Solids Content and Activity of raw and pretreated Amberlyst 15...... 123 Table 4.02 GC and Integrator Settings...... 124 Table 4.03 Calibrations for Analysis of TBA in Water/EG by FID...... 125 Table 4.04 Calibrations for Analysis of Ethylene Glycol mono-tert-Butyl Ether...... 125 Table 4.05 Calibration for Dissolved Isobutylene Analysis by FID...... 126 Table 4.06 Kinetic Study Experimental Range and Design...... 133 Table 4.07 Enthalpy and Entropy of Isobutylene Solubility...... 139 Table 4.08 Calculation of Internal Mass Transfer Parameters According to Equations (4.53) to (4.55)...... 146 Table 4.09 Enthalpy and Entropy of adsorption for Models M5 and M7...... 162 Table 4.10 Heterogeneous Kinetic Models Development and initial Parameter Estimates...... 163 Table 4.11 Results of Heterogenous Kinetic Model Further Evaluation...... 166 Table 5.01 Porosity and Static Holdup of Bale and Raschig Packing...... 188 Table 5.02 Fraction of Bale packing porosity occupied by liquid in the pre-loading regime and close to the flooding regime...... 197 Table 5.03 The corresponding flowrates of water and steam (saturated 373 K) to the loading ReG and ReL obtained and the ideal reboiler duty required to achieve the steam flowrates...... 197 Table 5.04 Bale packing Pressure Drop, ΔP/L (kPa.m-1)...... 204 Table 5.05 NORM and Sj values of Savitzky-Golay polynomial smoothing...... 206 Table 5.06 Basic Results of the Analysis of Gas Phase RTD curves for Raschig Ring Packing...... 210 Table 5.07 Basic Results of the Analysis of Gas Phase RTD curves for Bale Packing...... 211 Table 5.08 Indication of extent of backmixing, Levenspiel (1972)...... 222 Table 5.09 Kodak DC4800Camera settings for liquid RTD determination...... 228 Table 5.10 Basic Results of the Analysis of Liquid Phase RTD curves for Raschig Ring Packing...... 234 Table 5.11 Basic Results of the Analysis of Liquid Phase RTD curves for Bale Packing...... 235 Table 5.12 Lack of variation of dependent variables in ReG...... 235 Table 6.01 Wilson coefficients for TBA, Water and EG, Liu et al. (1993)...... 256 Table 6.02 Wilson coefficients for TBA and Water, Yang and Wang (2002)...... 256 Table 6.03 Molar Volume of TBA, Water and EG coefficients of Equation (6.03)...... 256 Table 6.04 UNIFAC Necessary Sub-Group Parameters...... 258 Table 6.05 UNIFAC Necessary Main Group Interaction Parameters...... 259 Table 6.06 Antoine Equation (6.18) Coefficients for TBA, Water EG and IB...... 260 Table 6.07 Determination of xrmin using Wilson and UNIFAC Models for the Extractive Distillation of Water and TBA...... 267 Table 6.08 GC and Integrator settings for analysis of water and TBA within water/TBA/Ethylene Glycol mixtures...... 272 Table 6.09 Calibration of area versus concentration for analysis of TBA and water...... 272 Table 6.10 Extractive Distillation of TBA and Water with Ethylene Glycol Main Results...... 280 Table 6.11 Extractive Distillation of TBA and Water with Ethylene Glycol, Column Temperature Profiles...... 281

xviii Table 6.12 Magnitude of reaction rates and isobutylene solubility required to attain the experimental flowrates of azeotrope used in this study, at 373 K, over 185 g Amberlyst 15, in an ideal CSTR...... 299 Table 7.01 GC and Integrator Settings for gas phase analysis of isobutylene...... 310 Table 7.02 Variable Ranges for the CFBR Recirculation Study...... 316 Table 7.03 Results of the Single Pass Water and Ethylene Glycol Containing Runs...... 319 Table 7.04 Physical Properties and Relations used to calculate mass transfer coefficients...... 323 Table 7.05 Mass Transfer Coefficients (kia) and Mass Transfer Step Resistances (Ri) Calculated...... 326 Table 7.06 Calculated Measures of the Experimental Design of the CFBR Recirculation Study...... 331 Table 7.07 TBA Reaction Rate, rTBA, ANOVA...... 332 Table 7.08 TBA Overall Effectiveness Factor, η0, ANOVA...... 333 Table 7.09 Rate of conversion and an estimate of conversion per pass...... 334 Table 7.10 MET Reaction Rate ANOVA...... 335 Table 7.11 MET Overall Effectiveness Factor ANOVA...... 336 Table 7.12 Selectivity Factor based upon Observed Reaction Rates ANOVA...... 336 Table 7.13 Total Observed Resistance for MET and TBA CFBR recirculation runs...... 339 Table 7.14 Experimental Design Factor Levels of CED Study Main Runs...... 347 Table 7.15 Solvent to water ratios obtained across the study...... 347 Table 7.16 Summary of Dependant and Independent CED Study Variables, Main (1 to 11) and Additional (12 to 17) Runs. For Stream Identification refer to Figure 7.11...... 349 Table 7.17 CED Column Temperature Profiles...... 350 Table 7.18 Isobutylene Balance on Main Runs under Zero Reflux...... 353 Table 7.19 Water Balance on Main Runs under Zero Reflux...... 353 Table 7.20 TBA Reaction Rate, rTBA.obs, ANOVA...... 354 Table 7.21 Intrinsic reaction rate for same temperature and concentrations as the average of the bale reactive zone...... 354 Table 8.01 Estimate of hole pitch, lp, required to achieve certain % Open Area...... 373. Table 8.02 Single Stage Tray study full factorial experimental design...... 377 Table 8.03 Single Stage Extended Capacity Study Experimental Variables...... 377 Table 8.04 BIC Initial Plate Design Study Results...... 379 Table 8.05 Physical Characteristics of the Basket Impeller Column...... 398 Table 8.06 Comparison of hL Predicted by Equation (6) and Correlations of Xu et al (1994) and Garcia and Fair (2002)...... 402 Table 8.07 The Effectiveness Factor with CW and dp...... 410 Table 8.08 Summary of Recent Kinetic Models for the Dehydration of TBA...... 412

xix 1. Introduction

Tertiary Butyl Alcohol (TBA) has multiple industrial applications. These include use as a denaturing agent; in the manufacture of flotation agents; developing flavours and perfumes; a solvent in paint removers and pharmaceuticals; as a dehydrating agent; and in the manufacture of methyl methacrylate, Clark (2001). Recent global trends in emissions legislation favour its use as a fuel additive for the purpose of boosting octane number of petrol, Stikkers (2002). Current Australian exhaust-gas control regulations prohibit the use of heavy metal-based additives in production, Australian Fuel Quality Standards Act 2000. TBA, produced by hydration of isobutylene, has blended properties akin to superior fuel oxygenates such as methyl tertiary butyl ether (MTBE) and ethyl tertiary butyl ether (ETBE) rather than other such as ethanol, Drogos and Dias (2001). The use of TBA for oxygenate has more favourable environmental implications than that of MTBE, the use of which has recently been scaled down.

TBA has above average fuel oxygenate properties, sounder environmental impacts and potentially cheaper feed stock than ethers. However, the economics of producing it by conventional means have traditionally been prohibitive, Piel and Thomas (1990). The hydration of isobutylene to TBA is a suitable candidate to share the benefits of the application of catalytic distillation technology as well exposed for the synthesis of methyl acetate, Towler (2000); and MTBE, Flato and Hoffmann (1992). The process intensification of catalytic distillation can offer capital cost savings and reduced operating costs through improved process yields and conversion. The hydration of isobutylene is equilibrium limited, Delion et al (1986), and suffers from product inhibition, Velo et al (1988). The countercurrent operation and reactant/product separation offered by catalytic distillation have been shown to improve conversion and alleviate such inhibition, Taylor and Krishna (2000). Given that the reaction is also exothermic, catalytic distillation technology can utilise this heat directly for the purpose of separation achieving a high degree of heat integration.

Chapter 1: Introduction 1 Following the first investigation of the application of catalytic distillation for the synthesis of TBA over Amberlyst 15 as conducted by Zhang et al (2001) two limiting system properties in the attainment of its fullest potential have been identified. Firstly, the existence of a particularly pressure insensitive homogeneous azeotrope occurring between the reactant (water) and product (TBA). Given the lack of variation in composition with pressure of this azeotrope, it can not be surpassed through column operating pressure manipulation. Further, given that it exists between reactant and product the azeotrope can not be reacted past in a reactive distillation system. Secondly, the potential benefits of catalytic distillation are limited by the lack of solubility of isobutylene in this matrix, which severely limits attainable reaction rates.

In order to overcome these above-mentioned hurdles a new and exciting process of catalytic extractive distillation is proposed. This novel process sees the combination of catalytic reaction in the presence of a solvent, with enhanced distillative separation through the alteration of relative volatilities. As part of the author’s honours thesis, Safinski (1999), a wide variety of solvents were screened. The glycol chemical series was identified as offering extractive distillation solvents, which satisfied all the requirements of a good entrainer, as well as rating well in terms of process sustainability. Additionally, the glycol series offers the opportunity to select an entrainer that enhances or reverses the natural tendency of separation of water and TBA. Upon further consideration of solvent inertness, ethylene glycol was considered to be least reactivate amongst the most effective solvents and was thus selected. Ethylene glycol affects the activities of the water and TBA in such a manner as to improve their relative volatility in the direction of their natural distillative separation tendency. It reduces the polarity of the liquid phase improving the solubility of isobutylene. The implementation of catalytic extractive distillation (CED) has the potential to combat the challenges facing the application of catalytic distillation for the synthesis of TBA.

The successful application of catalytic distillation technology requires the choice of suitable catalytic hardware, which allows for countercurrent operation without loss of capacity to excessive pressure drop. The effectiveness of the hardware for a given application is determined through its ability to accommodate for both the requirements of separation efficiency and catalytic activity. Bale packing was initially chosen for the

Chapter 1: Introduction 2 purpose of this study as it is an industrially established form of catalytic distillation hardware and it has been previously applied to a similar reaction (tertiary amyl alcohol) in the presence of high concentrations of solvent, Gonzales (1997). The basic capacity, Subawalla et al (1997); and mass transfer Zheng et al (1992), over Bale packing have been previously determined. However, there is a paucity of information regarding the underlying mixing characteristics of this packing. Characterisation of mixing will allow for better understanding of reactor performance and lead to improved catalytic distillation reactor design.

The specific requirements of a growing number of new catalytic distillation potential applications are not easily achieved with available catalytic distillation hardware, Krishna (2002). Further, the parameters defining the existing packing are difficult to vary effectively and often require complete column reconstruction and packing re- fabrication. As part of this study a new form a catalytic distillation reactor design was introduced. The new Basket Impeller Column (BIC) combines the high slip velocities and good liquid renewal properties of a stirred basket reactor with the good mass transfer characteristics and simple design of a dualflow column. This type of reactor design will allow for rapid process development and may be ideal for specialist applications where poor mass transfer substantially limits reaction rates.

The primary objectives of this work are to demonstrate the feasibility of catalytic extractive distillation for the synthesis of tertiary butyl alcohol employing the solvent ethylene glycol and to accommodate the requirements of this process with suitable catalytic distillation reactor design.

In summary the specific goals and aims of this project are to:

1. To characterise the solubility of isobutylene as a function of temperature and concentration across the complete matrix of water/TBA/ethylene glycol. 2. Obtain intrinsic kinetics for isobutylene hydration in the presence of the solvent ethylene glycol at the conditions of the proposed catalytic distillation column. 3. Evaluate the inertness of ethylene glycol within the proposed system.

Chapter 1: Introduction 3 4. Characterise the requirements of the separation present within the solvent containing process. 5. Characterise the reaction side requirements and extent of mass transport limitation particularly on the process limiting component isobutylene. 6. Research, develop, fabricate and validate a suitable catalytic distillation reactor design, which reaches an operable compromise between the separation and reaction requirements.

In the next chapter a literature review is conducted of topics relevant to the fulfilment of the above stated aims. In the chapter following that the findings of the literature review are used to shape the proposed process and determine the additional information necessary for its further development and assessment. The pertinent questions raised by the proposed process are formulated into a list of specific tasks. These form the basis of subsequent chapters.

1.1 Literature Cited

Clark, J. (2001). tert-Butyl Alcohol: Chemical Properties, Production and Use, Fate and Transport, Toxicology, and Detection in Groundwater and Regulatory Standards. Oxygenates in Gasoline. A. F. Diaz and D. L. Drogos, ACS Books. Delion, A., B. Torck and M. Hellin “Equilibrium Constant for the Liquid Phase Hydration of isobutylene over Ion-Exchange resins.” Ind. Eng. Chem. Process Des. Dev. 25: 889-893. (1986). Drogos, D. L. and A. F. Diaz (2001). Physical properties of fuel oxygenates and additives. Oxygenates in Gasoline. A. F. Diaz and D. L. Drogos, ACS Books. Flato, J. and U. Hoffmann “Development and Start-up of a Fixed Bed Reaction Column for Manufacturing Antiknock Enhancer MTBE.” Chem. Eng. Tech. 15: 193- 201. (1992). Gonzales, J. C., H. Subaealla and J. R. Fair “Preparation of tert-Amyl Alcohol in a Reactive Distillation Column. 2. Experimental Demonstration and Simulation of Column Characteristics.” Ind. Eng. Chem. Res. 36: 3845-3853. (1997). Krishna, R. “Reactive separations: More ways to skin a cat.” Chem. Eng. Sci. 57: 1491- 1504. (2002).

Chapter 1: Introduction 4 Piel, W. J. and R. X. Thomas “Oxygenates for reformulated gasoline.” Processing. (1990). Safinski, T., 2-Methyl-1,3-propanediol as a cosolvent towards the catalytic distillation of tert-butyl alcohol, Honours, UNSW, Sydney, Australia, 1999 Stikkers, D. E. “Octane and the environment.” The Science of the Environment 299: 37- 56. (2002). Subawalla, H., J. C. Gonzalez, A. F. Seibert and F. J. R. “Capacity and Efficiency of Reactive Distillation Bale Packing: Modeling and Experimental Validation.” Ind. Eng. Chem. Res. 36: 3821-3832. (1997). Taylor, R. and R. Krishna “Modelling Reactive Distillation.” Chemical Engineeering Science 55: 5183-5229. (2000). Towler, G. P. and S. J. Frey (2000). Reactive Distillation. Reactive Separation Processes. S. Kulprathipanja. Philadelphia, Taylor and Francis: Chapter 2. Velo, E., L. Puigjaner and F. Recasens “Inhibition by Product in the Liquid Phase Hydration of isobutene to tert-Butyl Alcohol: Kinetics and Equilibrium Studies.” Ind. Eng. Chem. Res. 27: 2224-2231. (1988). Zhang, C., Isobutene hydration in three phase reactors, PhD, UNSW, Sydney, Australia, 2001 Zheng, Y. and X. Xu “Study on Catalytic Distillation Processes: Part I: Mass Transfer Characteristics in Catalyst Bed within Column.” Trans. I Chem. E. 70: 459. (1992).

Chapter 1: Introduction 5 2. Literature Review

The objective of this chapter is to provide sufficient contextual information on the application, technology and design methodology for catalytic distillation to allow for the successful design of a effective TBA synthesis process. This literature review will allow for effective experimental designs and provide a basis for the interpretation of experimental results.

This literature review is compiled as such: an elementary perspective on the possible applications of TBA as fuels oxygenate is presented. Secondly the current methods of production and research into improved methods are reviewed. In this regard recent progressive contributions to the understanding of the heterogeneous kinetics of isobutylene hydration to TBA are detailed. Extractive distillation and catalytic distillation, which form the basis of the novel proposed process of catalytic extractive distillation are reviewed through the exploration of relevant applications and topics of current research. The emerging alternative applications of catalytic extractive distillation are considered. Finally, countercurrent fixed bed reactors are addressed including their possible use as an evaluation tool for the reactive section of the current process being considered.

2.1 TBA For The Purpose Of Petroleum Blend Additive

Recent global and local trends in emissions legislation may favour the use of tertiary butyl alcohol (TBA) as a fuel additive for the purpose of boosting octane number of petrol and this purpose was investigated and discussed in brief.

2.1-1 Octane Number

When petrol is burned in an internal combustion engine there is a tendency for it to burn unevenly depending on its composition. This is caused by ignition of the mixture before the piston is in the proper position. Such non-constant combustion can cause a knocking

Chapter 2: Literature Review 6 noise in the engine and is thus referred to as knocking. In order to quantify a petrol blend’s propensity for knocking a scale called Octane number was developed. Based on the combustion properties of organics known at that time 100 was assigned to isooctane and 0 to n-heptane. Mixtures of the two formed a scale in terms of the percentage of isooctane by volume. Petrol blends were then compared to these mixtures in terms of knocking and thus rated in terms of Octane number, Chenier . Several scales are used in the refinery industry today and vary in terms of experimental conditions, Stikkers (2002): 1. Research Octane Number, (RON), reflects performance at 600 rpm, 125 oF (51.7 oC) and low speed. It is used in Australia as the standard scale by refineries and petrol stations. 2. Motor Octane Number, (MON), reflects performance at 900 rpm, 300 oF (148.9 oC) and high speed. 3. Anti-Knocking Index (AKI) or RON-MON Average, this is the arithmetic average of the above scales, (RON+MON)/2. It is used by petrol stations in the USA and is commonly quoted in general literature.

The octane number varies with organic properties as follows: 1. The octane number increases as the amount of branching or number of rings increases. Eg. RON methylcyclohexane 75 / RON n-heptane 0. 2. The octane number increases as the number of multiple bonds increases. Eg RON 120 / RON methylhexane 75.

The order of decreasing Octane number for a given number is thus: aromatics, and > cyclic and branched alkanes > straight-chain alkanes. As combustion is a free radical process, the compounds which form more stable free radicals burn more easily and smoothly. Hence cyclic and branched compounds perform better than straight ones.

High octane reduces pollution in terms of carbon dioxide as it increases fuel economy through the ability to use higher compression ratios. Hence setting a minimum RON is necessary for engine efficiency and pollution control. Thus a major governmental tool

Chapter 2: Literature Review 7 in terms of pollution control is to set a minimum RON and regulate the maximum and minimum of allowed components.

2.1-2 Fuel Additives, Oxygenates and Others, History of Developments

Much research and effort has been given to finding means of increasing octane number without altering the hydrocarbon makeup of petrol by adding small amounts of components which would vary the octane number significantly. This approach has been considered to be much more economic than calling upon the existing refinery processes such as and polymerisation The manipulation of Octane number its history, development, legislation and environmental impacts has been reviewed from a US perspective by Stikkers (2002).

The first fuel additive developed was tetraethyllead, PbEt4, (TEL). During combustion the TEL formed lead dioxide, which complexed with and assisted carbon- carbon and carbon- bond breaking. The lead dioxide had to be removed from the engine and thus TEL was sold as a mixture called ethyl fluid containing halogenated hydrocarbons, which reacted with the lead dioxide and produced volatile compounds. Once released to the atmosphere elemental lead and halogen resulted through the action of atmospheric chemistry. These are obviously serious pollutants. Further the lead deactivated catalytic converters which had been developed to assist in controlling tail pipe emissions. In 1971 the US EPA issued the first regulation to stepwise reduce TEL. The final ban was passed in 1990 and the final date set to 1st January 1996.

In order to combat the loss of TEL refineries began to increase the production of high octane aromatic hydrocarbons by catalytic reforming. In the early 1980s the refineries used compounds such as and toluene to a large extent. The use of organometallic was also reconsidered based on earlier work with methylcyclopentadienyl manganese tricarbonyl (MMT), which was later used by Canada. The refineries next turned to high-octane alcohols and ethers and these additives came to be known as oxygenates as they contain an oxygen . Since these compounds include oxygen they improve combustion by making the fuel/air mixture leaner. Especially in older cars this enleanment effect mitigates incomplete combustion

Chapter 2: Literature Review 8 and thus reduces the emission of CO and unburnt hydrocarbons. In August 1989 ARCO announced a new fuel blend which incorporated MTBE and introduced the concept of reformulated gasoline (RFG). The RFG policies, which then took off, focused on the reduction of aromatics

While the fuel oxygenates were doing well in terms of atmospheric pollution control, it soon became apparent that oxygenates such as MTBE presented a threat to water supplies. The reasons for this problem are MTBE’s relatively high water solubility, objectionable odour and taste at very low concentrations and slow biodegradability. The MTBE saga came to a head when the city of Santa Monica California was forced to shut down 7 out of 11 drinking water wells, in 1995. Ultimately, California Governor at the time, Gray Davis signed an executive order in March 1999 directing state regulations to eliminate MTBE from gasoline by the end of 2002. In late 1999 a US EPA appointed commission on the threat of MTBE released it’s findings calling for a substantial reduction in MTBE use at a national level. In 2000 the possible carcinogenicity of MTBE was raised.

The use of oxygenates is currently being decided upon. In their favour is the fact that small quantities of oxygenates contribute significant amounts of octane to petrol blends. There is much protection and support for the use of ethanol worldwide. Ethanol carries it’s own range of associated problems. It increases the vapour pressure of the petrol blend and separates out to any water contaminating the petrol. It also has a tendency to increase NOx. Further the US EPA has upgraded acetaldehyde, a primary product of ethanol combustion, to likely human carcinogen.

Another issue regarding octane and refineries is the growing need to reduce the levels of sulfur in the petrol in order to boost catalytic converter performance and lower emissions further. Of the current methods of desulphurisation used in the refineries the most cost effective ones also reduce octane, Stikkers (2002).

2.1-3 Fuel Oxygenates, the Global Situation

In Europe, MTBE is currently commonly used at concentrations around 1.6%, and apart from Denmark which extracts a large proportion of its drinking water from underground

Chapter 2: Literature Review 9 resources, there appear to be no plans to prohibit MTBE under the new Euro fuel standards. MTBE is, however, being monitored closely for occurence in ground water, Schmidt et al (2001). The USA is currently phasing out MTBE, with California leading the way.

Table 2.01 Maximum concentration of compounds under EU and US legislation Property EU EU EU EU USA USA EURO 1 EURO 2 EURO 3 EURO 4 RFG CA2RFG Effective 1985 1995 2000 2005 Aromatics v % - - 42 35 25 30 Benzene vol % - 5 1 1 1 1.2 Oxygen mol % 2.5 – 3.7 2.5 – 3.7 2.7 2.7 2.7-3.5 1.8-2.7 Ethers vol % 10 – 15 10 - 15 15 15 - - Sulfur mg/kg - 500 150 50 - 80 Lead g/L - 0.013 0.005 0.005 - 0.013

2.1-4 The Australian Situation

Currently the specification of petrol in Australia falls under the newly implemented Fuel Quality Standards Act 2000. Even though Australia is lagging behind the US and EU in terms of emissions standards, the Act is a positive development to close the gap. Australia is following the EU model which is set out in Standards EURO 3 and EURO 4, see Table 2.01. EURO 3 was implemented in 2000 and EURO 4 is to be implemented in 2005. Australia hopes to have adopted EURO 3 by 2005/6 for petrol and EURO 4 by 2006/7 for diesel. All common fuel oxygenates other than ethanol are currently prohibited from use in petrol blends. The values permitted are such as to allow for contamination, 0.5 wt%, however an application can be made for Approval to vary the petrol standard under the Act. In adopting the European standards the use of oxygenates will certainly come into consideration within the next decade. Further with MTBE falling out of favour and with its favourable properties TBA may be considered as a potential fuel additive.

2.1-5 TBA as Fuel Oxygenates

TBA possesses many of the properties required of a good fuel oxygenates. It is an attractive alternative to MTBE. The feedstock of water provides for a cheaper reactant than methanol and ethanol for reaction with isobutylene. It can also be used as mixture with methanol or ethanol. The main disadvantages of using simple alcohols, such as

Chapter 2: Literature Review 10 ethanol and methanol relative to ethers such as MTBE, ETBE and TAME according to Stikkers (2002) were found to be: 1. Alcohols were difficult to mix with petrol because of their tendency to separate from the fuel blend in the presence of water. Which means that they are difficult to blend at the refinery and transport using existing infrastructure. 2. Alcohol-based additives can significantly increase the vapour pressure of the fuel mixture. This can increase the evaporative emissions, calling for new expensive control measures.

A comparison of oxygenate properties in Table 2.02 shows that these disadvantages typically associated with alcohols are substantially reduced in the case of TBA.

Table 2.02 Physical Properties of Fuel Oxygenates, Drogos and Dias (2001) Property MTBE TAME ETBE EtOH MeOH TBA mp oC -108.6 - -94 -114 -97.6 25.4 bp oC 55.2 86.3 72.2 78 64.6 82.4 ρ (g/cm3) @ 25 oC 0.7352 0.7656 0.7353 0.7900 0.7870 0.7750 Pvap mmHg @ 25oC 245 68.3 152 50 126 41 o log Koc @ 25 C 1.035- 1.27-2.2 0.95-2.2 0.2-1.21 0.44-0.92 0.37-1.57 1.91 o Log Kow @ 25 C 1.20 1.55 1.74 -0.31 -0.77 0.35 blend (wt % O2 / 2.7/15.0 2.7/16.6 2.7/17.1 2.7/7.5 - 2.7/11.8 vol % O2) RON 117 112 118 130 133 109 MON 101 98 101 96 105 100 Blend RVP (psi) 8 2.5 4 18 40 9 Neat RVP (psi) 7.8 2.5 4.0 2.3 - 1.7

The blended vapour pressure is much the same as that of MTBE and thus the emission problem associated with alcohols would not exist. Further the partition coefficient of

TBA between octane and water, Kow is not as unfavourable as that of MeOH and EtOH in terms of water pollution and not as high as that of the ethers in terms of adsorption in body fats. The organic-carbon partitioning coefficient Koc is defined as the ratio of adsorbed chemical on organic carbon to concentration in water. The low Koc of TBA means that it is not likely to attach to organic matter in soil or be removed from aqueous solution by granulated organic carbon. The values of Koc are higher for ethers, but lower for the other alcohols. The other alcohols however would have a much higher limit of toxicity. The RON and the MON of TBA compare well with both ethers and other

Chapter 2: Literature Review 11 alcohols. The main disadvantage holding back TBA, is the economics of producing it by conventional means, Piel and Thomas (1990).

TBA has the potential to be a ground water contaminant much like MTBE. It is highly mobile and is not adsorbed by organic carbon in soils to any appreciable extent. It is also not easily biodegraded, due to the resistance of tertiary and quaternary to microbial attack. It is not highly toxic to humans or animals. It has a lower toxicity to aquatic life than other oxygenates such as MTBE. The lethal concentration resulting in 50 % population death was 2450 to 6410 mg/L compared to that of MTBE of 672 mg/L. It’s carcinogenicity is low. Its genotoxicity is very low. US ground water limits for TBA range from 12 to 100 μg/L, Clark (2001).

While microbial attack is limited it has been shown to occur and to be favoured by aerobic conditions, Bradley et al (2002) and Deeb et al (2001). The cellular yield of the organism using TBA as carbon source was low. The presence of more easily biodegradable compounds strongly inhibits it’s degradation. However the bioremedation of contaminated soil is seen to be possible for both MTBE and TBA, Deeb et al (2001).

The organic emissions, volatile organic , VOCs, can compete with ozone in the atmosphere in the presence of light to react with NO thus resulting in an increase in localised ozone and the production of smog. Some research has been conducted to evaluate the reactivity of the fuel oxygenates compared to the other petrol constituents emitted such as alkanes, aromatics alkenes and carboyls. Amongst the fuel oxygenates themselves those containing ethyl groups such as ethanol, ETBE, TAME are more reactive then those containing methyl or tert-butyl groups, such as MeOH, MTBE and TBA. Both on a carbon and molecule basis TBA has the lowest reactivity. Thus in terms of unburnt fuel oxygenates as VOCs, TBA would have the lowest smog contribution, Bowman and Seinfeld (1995).

Chapter 2: Literature Review 12 2.2 Industrial Production of TBA

Industrially TBA has been produced both as a byproduct of the propylene oxide process and a product of both homogeneously and heterogeneously catalysed hydration of isobutylene.

2.2-1 TBA as a By-product of the Halcon Propylene Oxide Process

Currently TBA is produced commercially as a byproduct in the Halcon Process for the manufacture of propylene oxide (1,2-epoxypropane), Chenier . Propylene oxide is itself a precursor for propylene glycol, glycol and various glycol ethers. Oxygen is first used to oxidise isobutane to t-butyl hydroperoxide (TBHP) and TBA over a molybdenum naphthenate catalyst at 90 oC and 450 psi as follows:

CAT 4 (CH3)3CH + 3 O2 ⎯⎯⎯→ 2 (CH3)3COOH + 2 (CH3)3COH (2.01)

In the next reaction step the TBHP is used to oxidise propylene to the oxide with a yield of 90 % relative to propylene:

CAT (CH3)3COOH + CH2=CHCH3 ⎯⎯⎯→ CH2OCHCH3 + (CH3)3COH (2.02)

It can be seen that TBA is produced in both steps. Approximately 2 kg of TBA are produced for every 1 kg of propylene oxide. Lyondell the world’s largest manufacturer of propylene oxide currently has two facilities using this process, as well as an alternative process which produces styrene as a byproduct, the Bayport Plant, in Pasadena Texas and the Channelview Complex, Channelview Texas.

2.2-2 Direct Hydration of Isobutylene

There is strong interest in the use of the direct hydration of Isobutylene (IB) to TBA, due to the relatively lower cost of the feedstock. The hydration of IB to TBA over an acid catalyst follows the general formula:

H+ (CH3)2CCH2 + H2O ←⎯⎯⎯⎯→ (CH3)3COH (2.03)

Chapter 2: Literature Review 13 TBA has classically been synthesised by homogeneous catalysis with sulfuric acid, Kroper (1969). Isobutylene reacted with sulfuric acid (45 wt%) to form the sulfate followed by hydrolysis. The use of sulfuric acid requires reconstitution of dilute acid by-product and the process is energy inefficient and corrosive. Further the extent of side reactions forming di and trimers of isobutylene is significant. The production of pure isobutylene by extraction with sulfuric acid (50 wt %) and dehydration of TBA was also developed, Martel et al (1965). British Hydrocarbon Chemicals commissioned a CFR designed plant in 1963, producing 100 ton/day of isobutylene. The reaction vessels and towers were largely constructed of carbon steel lined with lead. Monel was used in areas of high velocity and high temperature.

The disadvantages of homogeneous catalysis have led to considerable interest in heterogeneous catalysis Gupta and Douglas (1967). Heterogenous catalysis over ion exchange resins has been explored and implemented, however there is a paucity of public information for the industrial processes developed. The Hüls Process for the purification of isobutylene from a mixed butylene stream incorporated the hydration of isobutylene over an ion exchange resin to TBA, Hüls (1983). The plant produced 50,000 metric tons/year of very pure TBA. The reactors employed were operated at staggered temperatures with multistage addition of water. The process was operated with hydrocarbon feed in excess and a large amount of DIB was also produced. Several reactors and multiple columns combined with recycle streams where employed. The use of ion exchange resin allowed the plant to be built of carbon steel.

The reactant isobutylene typically is available as part of refinery gas stream. Steam of naphtha produces ethylene, propylene and appreciable amounts of a C4 fraction. The C4 fraction includes , isobutane, 1-, cis and trans-2-butene, isobutylene and . A typical composition of the C4 stream is given in Table 2.03. The butadiene is most commonly used in the and rubber industry, while isobutylene is typically used to enrich the octane rating of petrol either through alkylation, polymerisation or reaction to oxygenates.

Chapter 2: Literature Review 14 Table 2.03 Composition of a typicalC4 Stream Component mol % n-butane 3 isobutane 1 isobutene 23 1-butene 14 2-butene 11 butadiene 47 other 1

Previously the butadiene was removed by cuprous ammonium acetate extraction, leaving a mixture of isobutylene and n-butene, Martel (1965). It is this stream that most studies in the synthesis of ethers, such as MTBE and ETBE, refer to. According to

Jayadeokar and Sharma (1993) on a commercial scale, isobutylene comes from the C4 fraction of the units with an isobutylene content ranging from 40 – 50 mol % after butadiene extraction or catalytic crackers with isobutylene content of 15 – 20 mol %. Good process design and optimisation requires a sound knowledge of reaction kinetics, chemical and physical equilibrium. These will be considered in the next section for the direct hydration of isobutylene to TBA.

2.3. Thermodynamics and Kinetics of the Hydration of Isobutylene

2.3-1 Thermodynamics of Isobutylene Hydration

Considered in this section are chemical equilibrium of reaction equilibrium constant and physical equilibrium of isobutylene solubility and miscibility. Delion et al (1986) proposed that the thermodynamic equilibrium constant K in temperature followed:

⎛⎞4264.7 ln K = ⎜⎟− 10.945 (2.04) ⎝⎠T

This correlation gives K constants of 9.57 at 50 oC, 3.11 at 80 oC and 1.62 at 100 oC. Thus the reaction is considerably equilibrium limited. The standard enthalpy of the reaction is –35.45 kJ/mol. Hence the reaction is moderately exothermic.

Chapter 2: Literature Review 15 The solubility of isobutylene in water was determined gravimetrically by Kazanskii et al (1959). Leung et al (1987a) extended the study to mixtures of water and TBA. Their gas chromatography based analysis was calibrated against the results of Kazanskii et al. (1959). Table 2.04 Solubility of isobutylene within mixtures of water and TBA T oC S ×1003 mol.L-1* S ×1003 mol.L-1 ** S ×1003 mol.L-1 ** TBA mol/L: 0 1.04 2.95 30 4.66 5.50 33.0 40 3.55 4.18 28.4 50 2.68 3.23 26.5 60 1.93 2.60 20.3 * Kazanskii (1959), **Leung (1987a)

Leung et al (1987a) suggested the following correlation for pure isobutylene at pressure of 101.3 kPa in the presence of TBA (0≤CTBA ≤ 2.95 mol/L) and over the temperature range 303 ≤ T ≤ 333 K:

⎛⎞A sBIB =+exp⎜⎟ ⎝⎠T 2 B =−16.975 + 2.101CCTBA + 0.259 TBA (2.05)

AC=−2322 662.4 TBA

Zhang, Adesina and Wainwright (UNSW(2002)) investigated the effect of TBA upon solubility of isobutylene within the primary solvent water in much the same manner as Leung et al (1987a). They obtained the following correlation for the ratio of solubility of isobutylene within the mixture to that within pure water:

03 sIB. mix −02 ⎛⎞−×29 10⎛⎞ − 19.96 =×6.155 10 exp⎜⎟exp⎜⎟ (2.06) sTCIB. W⎝⎠8.314 ⎝⎠ TBA which was intended to be used with their experimental pure water Henry’s constant of 03 -3 -1 HIB.W = 20.7×10 kPa.m .kmol .

2.3-2 Homogeneous Kinetics of Isobutylene Hydration

The hydration of isobutylene is specifically acid catalysed and so is catalysed only by the solvated hydrogen ion. Taft (1952) first proposed the mechanism of the reaction.

Chapter 2: Literature Review 16 The mechanism involved the reversible protonation of the olefin yielding a π-complex Equation 2.07, this was further transformed into a carbonium ion Equation 2.08 the rate determining step. This carbonium ion binds with water forming TBA releasing the proton Equation 2.09.

Fast CH C=CH + H O++ ⎯⎯→ CH CH =CH + H O (2.07) ()32222←⎯⎯ ()3 22

Slow CH CH++ =CH ⎯⎯⎯→ CHC (2.08) ()3223←⎯ ⎯⎯ () 3

Fast CH C++ + H O ⎯⎯→ CH COH + H (2.09) ()3233←⎯⎯ () 3 This mechanism has strong proponents in Lutsyk et al (1998), who have extended its validity to concentrated acids, using the Hammett acidity function. It has also been challenged by Baliga and Whalley (1964). They argued that such a mechanism should be accompanied by a volume change and showed experimentally that this did not occur. They further noted that the mechanism should give positive entropy of activation, while that calculated by Taft was negative.

Gehlawat and Sharma (1968) investigated the rates of absorption of isobutylene in sulphuric acid solutions of concentrations of industrial significance of 50 to 70 wt/wt. Using Dankwerts’ surface renewal theory the rate of absorption R mol/cm2.s was defined:

2 R = sDkkIB+ L (2.10)

3 2 where sIB was solubility of IB in the electrolyte (mol/cm ), D diffusivity (cm /s), k pseudo first order reaction rate constant and kL physical absorption coefficient without reaction. The parameters D, kL and sIB were evaluated experimentally in specific studies. Thus upon measuring rates of absorption k could be determined and was correlated as:

−05 k =×4.1 10 exp 2[ H24 SO ] (2.11)

The apparent energy of activation was found to be 54.39 kJ/mol.

Chapter 2: Literature Review 17 2.3-3 Heterogeneous Kinetics of Isobutylene Hydration

Currently the popularised catalyst for the hydration of isobutylene is Amberlyst 15, a sulfonic acid macroreticular ion exchange resin manufactured by Rohm and Haas, Velo et al (1988). The choice of Amberlyst 15 lies in its strong affinity for water, higher internal diffusivities than other resins, good stability and appropriate temperature range, as well as good commercial availability and a sound base of characterisation data, Leung et al (1986). Amberlyst 15 is a styrene-divinylbenzene copolymer with sulfonic functional groups. It can be obtained dry with less than 1% moisture or wet with 50 – 55 % moisture. The nearly spherical beads of the catalyst are agglomerates of gel-type microparticles surrounded by a macroporous matrix. The microparticles have been estimated as having a diameter of 0.1 μm, Dooley et al (1982), other properties are listed in Table 2.05. Different rates of diffusion can exist in each of these regions of the macroporous catalyst, Ihm, Chung and Park (1988).

Table 2.05 Properties of Amberlyst 15 dry, Rohm and Haas (R&H) Property Value Source Max operating Temp (K) 393 R&H Ionic Form Hydrogen R&H Capacity (eq mol H+.kg-3) 4.7 R&H 4.9 Rehfinger (1990) Particle Size dry (μm) 350-1200 R&H Surface Area (m2.kg-1) 50 × 103 R&H 49.4× 103 Rehfinger (1990) Porosity 0.36 R&H 0.35 Leung (1986) Avg. pore diameter, (nm) 24 R&H 26 Leung (1986) Hydrated Density (kg.m-3) 927 Leung (1986) Skeletal Density (kg.m-3) 1426 Leung (1986) Swelling in Water (%) 60 - 70 Ihm (1988)

The hydration of isobutylene is said to be catalysed by the hydrated protons rather than by H-SO3 groups. The mechanism for this reaction in the presence of excess H2O has been proposed as follows

+ k1 + (CH3)2CCH2 + H3O ←⎯⎯⎯⎯→ (CH3)3C (2.12) k−1

+ k2 + (CH3)2C + 2H2O ←⎯⎯⎯⎯→ (CH3)3COH + H2O (2.13) k-2

Chapter 2: Literature Review 18 It involves the formation of tert-butyl cation, which may revert to isobutylene or be hydrated to the alcohol. Additionally this cation can potentially react with isobutylene to form diisobutylene (DIB) or react with a TBA molecule to form di-tert-butyl ether. However according to a healthy body of previous research, these possible side reactions do not take place to any appreciable extent over heterogeneous catalysts in the presence of water, Velo et al (1990). In the absence of water IB undergoes oligomerisation readily to di- tr- isobutylene over A-15 at a moderate 60 oC and100 kPa, Alcantara et al (2000). The formation of the sec-butyl cation instead of the tert-butyl cation does not occur. The reaction follows very strongly the Markovnic rule wherein the addition of H- X to an the hydrogen atom adds to the carbon of the with the greater number of hydrogen atoms.

The earliest application of ion exchange resin catalysis for the hydration of isobutylene was that of Gupta and Douglas (1967). They employed the hydrogen form of Dowex 50W, monosulfonated polystyrene crosslinked with 8% divinyl benzene. This is a gel type resin and it has been since been demonstrated that intraparticle diffusivities are much higher than for macroporous resins such as Amberlyst 15, Leung et al (1986). Two particle sizes of the gel resin were used to facilitate the determination of diffusion and reaction parameters. The pseudo-first order kinetics in isobutylene was determined in a pressurised system consisting of two liquid phases. They showed that the reaction rate was independent of the volumetric ratio of the reactants. A strong argument was made for the surface diffusion of isobutylene as De was found to decrease with increasing temperature. Gupta and Douglas proposed that for their range of experimentation (74 < T < 97 oC) the concentration of isobutylene at the surface of the -03 catalyst was almost constant at CiB.S = 0.0172 ×10 mol/mL swollen resin.

The approach of Gupta and Douglas was extended to Amberlyst 15 and Amberlyst XN- 1010 by Ihm, Chung and Park (1988). They divided the rate of reaction observed between that occurring on the surface of the microparticles and within these microparticles, which make up the macroreticular resin. The reaction within Amberlyst XN-1010 was assumed to occur on the exterior of the microparticles as the high extent of crosslinking of the resin 70-75% prevented swelling. They found that the internal

Chapter 2: Literature Review 19 sites were more active and the effectiveness factor for the gel microporous particles was equal to unity, η = 1.

Table 2.06 Internal mass transfer parameters: De, φ and η, for isobutylene hydration over Amberlyst 15, Leung et al (1986)

dp (μm) T (K) η φ De ×1010 m2/s 450 303 0.78 0.73 3.7 313 0.69 0.95 5.1 323 0.62 1.15 8.0 333 0.52 1.51 11.0 1040 303 0.48 1.67 3.7 313 0.38 2.20 5.1 323 0.33 2.65 8.0 333 0.26 3.48 11.0

Leung et al (1986) used a recirculation, differential packed bed reactor in order to obtain mass transfer free kinetics for the evaluation of isobutylene hydration within a trickle bed reactor Leung et al (1987b). They used particles of diameters dp = 450 and 1040 μm in order to characterise the effectiveness factor with temperature (303 – 333 K) and found that it decreased with temperature and decreased with particles size, Table 2.06. They found that the intrinsic reaction rate was pseudo first order in isobutylene:

−−11 ⎛⎞−8463.10 rmolgsTBA (). .= 53440786.14exp⎜⎟CIB (2.14) ⎝⎠T and that η = 1 in the microparticles collaborating the findings of Ihm et al.

Velo et al (1988) used a recirculation, differential packed bed reactor in the same manner as Leung et al (1986) but extended the range of study to include product addition. By adding TBA to their feed up to the concentration of 3 mol/L they were able to demonstrate that it has an inhibiting effect upon reaction rate. This inhibition effect referred to the observation that while increased TBA concentration promotes isobutylene solubility there was no proportional gain in reaction rate. Further that as equilibrium is approached the reverse reaction begins to slow down the apparent rate of the forward reaction. The model proposed by Velo et al represented 80% of the observed data with a 10 % error. It was defined in the range 0 ≤ [TBA] ≤ 3 mol/L, with

Chapter 2: Literature Review 20 liquid saturated with IB with partial pressures up to 250 kPa, over the temperature range 30 to 60 oC. They could not distinguish between exponents n =1 or 2 and thus quote both models:

⎛⎞CA kCC⎜⎟BW− ⎝⎠KC r = n (2.15) ()1 + KAA C

where CB, CW and CA are the concentrations of IB, water and TBA respectively. The constant are given as a Arrhenius function of temperature:

⎛⎞8844 k = exp⎜⎟ 15.03 − (2.16) ⎝⎠T

The overall equilibrium constant KC is defined as:

⎛⎞3160 KC =− exp⎜⎟ 6.78 (2.17) ⎝⎠T

The TBA inhibition constant KA depends on the order of the denominator, n. For n = 1: ⎛⎞8540 KA =− exp⎜⎟ 26.6 (2.18) ⎝⎠T and for n = 2: ⎛⎞6602 KA =− exp⎜⎟ 19.49 (2.19) ⎝⎠T

The temperature range of their model is limited (30 to 60 oC) and would need to be extended for use with catalytic distillation work. Velo observed that the reaction rate was first order with respect to IB.

Zhang, Adesina and Wainwright (UNSW(2003)) have confirmed this first order dependency in isobutylene in both gas and liquid concentrations obtaining at 343 K the following pseudo first order rate expressions:

Chapter 2: Literature Review 21

−07 rPTBA =×1.34 10 IB (2.20) and

−03 rCTBA =×7.26 10 IB (2.21)

With additional measurements at 343 K of reverse reaction rate they were able to propose the following simple and yet versatile model:

−03 ⎛⎞CTBA rCTBA=×7.0 10 ⎜⎟ IB − (2.22) ⎝⎠26

Their equilibrium constant compared well with the work of Delion et al (1986). For a wider range of temperature (323 – 353 K) they obtained the following:

−−11 +06 ⎛⎞−8323 rmolgsTBA (). .=× 2.07 10 exp⎜⎟CIB (2.23) ⎝⎠T

The apparent activation energies observed for the studies mention were reported in Table 2.07. It can be seen that they are generally lower than those reported for homogeneous catalysis.

Table 2.07 Activation Energies reported for heterogeneous catalysis over ion exchange resins. Study Catalyst apparent E kJ/mol Gupta and Douglas Dowex D50 95 (1967) Velo et al (1988) Amberlyst 15 73.7 Leung et al (1987b) Amberlyst 15 67 Ihm et al (1988) Amberlyst 15 61.5 Zhang et al (2003) Amberlyst 15 69

The inhibition effect described by Velo et al. (1988) has also been observed by Petrus et al (1984) in their study of the hydration of . They found that the value of their rate constant decreased by a factor of 2 as the 2-propanol concentration was increased

Chapter 2: Literature Review 22 from 0 to 10 wt %. They proposed that the propanol preferentially held onto the propene in solution preventing it from forming a carbonium ion. Their power law reaction rate was related to the amount of propene forming the carbonium ion and hence incorporated the inhibition effect.

Velo et al (1990) proposed in a later study that the strength of the observed inhibition effect might be due to a change in the rate of effective diffusivity, De, of isobutylene within the resin as TBA is formed. They found that indeed the intraparticle diffusivity decreased with increased TBA, but concluded that more work was required before a complete model was developed. They correlated the findings with the following:

⎛⎞3690 DCeA=−−exp⎜⎟ 9.52 0.551 − (2.24) ⎝⎠T

o for CA up to 2.45 mol/L, CB below 0.01 mol/L and 30 to 60 C. The negative temperature coefficient, it was suggested is possibly due to surface diffusion. The argument for surface diffusion of isobutylene has gained increasing support. Isobutylene having a high affinity for the polymer backbone of the catalyst compared to the aqueous solution. The formation of TBA begins to upset this and the isobutylene enters the liquid solvent where its transport is slower than upon the surface. Leung et al (1986) and Velo et al (1990) found that the apparent tortuosity factor decreased with increased temperature and that it increased with TBA. This was typical behaviour of solute/catalyst systems, which exhibit surface diffusion. The analysis of intraparticle diffusion of Gupta and Douglas (1967), over Dowex D50 had also suggested surface diffusion of isobutylene.

Aiouache and Goto (2003) measured and analysed the kinetics and equilibria of the etherification of tert-amyl alcohol with ethanol in terms of swelling equilibria and the Florey-Huggins model. They found that the water produced was strongly retained by the resin, during the course of the reaction.

Given that the hydration of isobutylene is an equilibrium reaction, further insight may be gained by considering the dehydration of TBA. The kinetics of the dehydration of

Chapter 2: Literature Review 23 TBA have been studied in their own right as there are several practical applications of this reaction. It has been identified as a potential catalytic distillation heat pump, which can employ the low value waste heat of a process. It is used to obtain IB from a TBA feed stock for the manufacture of MTBE and it is a useful relatively simple catalytic distillation test system.

Heath and Gates (1972) observed two significant regions in the effects of water upon the reaction rate of TBA dehydration over initially dry resins: 1. initially the water produced accelerates the reaction as it swells the resin allowing for improved mass transport to the interior. 2. further production of water results in product inhibition due to competition for catalytic sites. The transition between region 1 and 2 was most noticeable for gel type resins and was insignificant for Amberlyst 15 a macroporous type resin. Hence Amberlyst 15 did not display a delay or induction period and was a much more effective catalyst. Gates and Rodriguez (1973) demonstrated that the strong inhibition of water present at the lowest concentrations in Amberlyst 15 catalysis was due to the conversion of -SO3H groups to hydrated protons. They found that the first order rate constant for reaction catalysed by

–SO3H groups at very low coverage by alcohol was 40 times the first order rate constant for reaction catalysed by hydrated protons at low alcohol concentrations.

Kato et al (1996) obtained the following rate expression over A-15 for use with their heat pump research:

⎛⎞2863 exp⎜⎟ 1.92 − CTBA T ⎡⎤+ ⎝⎠ (2.25) rTBA ⎣⎦mol/(s.(molH )) = ⎡⎤⎛⎞⎛⎞⎛⎞⎛⎞141022 10245 1+− exp 55.795 +CC +− exp 36.938 + ⎢⎥⎜⎟⎜⎟⎜⎟⎜⎟HO2 IB ⎣⎦⎝⎠⎝⎠⎝⎠⎝⎠TT

Interestingly the kinetic model given by Equation suggested isobutylene adsorption in the reverse reaction. This indicates that it is possible that a different mechanism operates here to that of the forward reaction as proposed by Velo et al (1988).

Chapter 2: Literature Review 24 Abella et al (1999) conducted a study on dehydration recently over A-15 and found that the wet and dry form of the catalyst gave different reaction rates. The prehydrated wet form inhibited the initial reaction rate and it is the form, which is most likely to represent reaction in a system containing any significant amounts of water. Similarly to Kato et al (1996) they obtained a model that contained a water adsorption term however they did not find adsorption of isobutylene or TBA:

⎛⎞17017 exp⎜⎟ 39.6 − CTBA T ⎡⎤+ ⎝⎠ rTBA ⎣⎦mol/(s.(molH )) = (2.26) ⎡ ⎛⎞⎛⎞6669 ⎤ 1+−+ exp 22.3 C ⎢ ⎜⎟⎜⎟HO2 ⎥ ⎣ ⎝⎠⎝⎠T ⎦

The use of TBA as a direct precursor to MTBE and ETBE has been also investigated but largely hampered by difficulties, which demonstrate some of the key properties of hydration of isobutylene over Amberlyst 15. Yang et al (2000) found that the initial presence of water inhibited their reaction by a factor of as much as 20 times. Further that the TBA reverted quickly to water, which eventually blocked the reaction. Their results show the strength of prehydration of Amberlyst 15 on influencing subsequent reaction and the high affinity of Amberlyst 15 towards water leading to inhibition in systems where water is an inert or not the desired reactant.

2.3-4 Heterogeneous Kinetics of Similar Systems

Systems considered similar were ones identified as following the generalised formula:

+ ” Hcat ” R’C(CH3)C=CH2 + R OH ←⎯⎯⎯⎯⎯⎯→ R’C(CH3)2OR (2.27)

where R typically represents the end groups -H, -CH3 and -CH2CH3. The major systems defined by Equation 2.25 are those producing: 1. tert-Amyl Alcohol 2. Methyl tert-Butyl Ether 3. Ethyl tert-Butyl Ether 4. Tert-Amyl Methyl Ether

Chapter 2: Literature Review 25 Their reactions are catalysed by strong acid ion exchange resins and are usually carried out in the presence of inert components. In the case of isoamylene it should be noted that there are two 2-methyl-1-butene and 2-methyl-2-butene involved, giving one product of either TAA or TAME. The kinetics of the first three systems outlined will be explored in greater depth in this section.

Gonzalez and Fair (1997) investigating isoamylene hydration to tert-amyl alcohol and reported inhibition by water. At water concentrations higher than 5 mol % the isoamylene was reported to have been virtually excluded from the resin catalyst. To maintain a single liquid phase acetone was employed as a solvent. The kinetic behaviour was modelled in component activities using a modified Langmuir- Hinshelwood model. In this model the adsorption of water on resin was assumed to follow an exponential distribution:

⎡ 14 ⎛⎞−−8359 3.64 17 ⎛11544 ⎞⎤ ⎢1.01×− 10 exp⎜⎟aaIC5 W 1.95× 10 exp⎜ ⎟aTAA ⎥ ⎣ ⎝⎠TT ⎝ ⎠⎦ rTAA = (2.28) 3.64 2 ()1+ 26.2aW

The solvent acetone was absent in the kinetic model of equation 2.26. In light of the proposed kinetics of Velo et al for isobutylene hydration (1988), it seems that with increasing molecular size olefine hydration switches from being product to reactant inhibited. Gonzalez and Fair (1997) reported 69.5 kJ/mol and Delion et al (1987) reported 90.4 kJ/mol for the hydration of isoamylene over Amberlyst 15.

The most extensively referenced kinetics for the reaction of isobutylene and methanol in the presence of 1-butene to give methyl tert-butyl ether MTBE, are those of Rehfinger and Hoffmann (1990). The magnitude of the initial reaction rates observed (xEther = 0), -04 -05 for the range xMeOH = 0.2 to 0.8 was 1 × 10 to 5 × 10 mol/g.s over Amberlyst 15. Reaction rate was favoured by higher isobutylene fraction, at the expense of poorer parallel reaction selectivity and the formation of diisobutylene. They developed the reaction rate model given by:

⎡ aaIB MTBE ⎤ rqk=−r ⎢ 2 ⎥ (2.29) ⎣aKaMtOH a MtOH ⎦

Chapter 2: Literature Review 26 where r represents the reaction rate per unit catalyst mass, q is the amount of acid groups on the resin per unit mass (Amberlyst 15 = 4.9 eq/kg), and a is the activity of a component. The reaction rate constant k was given by:

12 ⎛⎞−11110 kr =×3.67 10 exp⎜⎟ (2.30) ⎝⎠T

The lumped equilibrium constant was given by:

KfTa = 284exp( ( )) (2.31) where

3 ⎛⎞⎛⎞11 T fT()=×−−1.49277 10⎜⎟⎜⎟ 77.4log + 0.5076() T − 298.15 ⎝⎠⎝⎠T 298.15 298.15 −×9.127 10−−−04()TTT 2 − 298.15 2 + 1.1065 × 10 06 () 3 − 298.15 3 − 6.2799 × 10 10 () 4 − 298.15 4 (2.32)

The apparent activation gained was 92.4 kJ/mol. This was adjusted for the enthalpies of adsorption giving 82.4 kJ/mol.

Fite et al (1998) studied the dependence of kinetic rates of the reaction of isobutylene with methanol on the swelling equilibria of the resin phase. They introduced a solubility parameter based upon Hildebrand theory. They recognised that the polarity of the fluid phase influences swelling, which in turn can influence the accessibility of active sites.

For the synthesis of ETBE by the etherification of isobutylene the kinetics have also been studied. Two sets of independently developed kinetics point towards ethanol inhibition and a mechanism involving three species in the rate-determining step. Two adsorbed ethanol species react with one free isobutylene molecule. Jensen and Datta (1996) proposed the following expression for the equilibrium constant in temperature: 4060.59 KT=+10.387 − 2.89055ln ETBE T (2.33) −+×−×0.01915TT 5.2859 10−−05 2 5.3298 10 08 T 3

Chapter 2: Literature Review 27

Jensen and Datta developed the following generalised kinetic model

2 ⎛⎞aETBE mkaEtOH⎜⎟ a IB − ⎝⎠KaETBE EtOH rETBE = 3 (2.34) ()1+ KaAEtOH

12 ⎛⎞−60.4 r =×7.418 10 exp⎜⎟ (2.35) ⎝⎠RT

1323.1 lnK =− 1.0707 + (2.36) A T

This generalised model could be simplified for EtOH > 4 mol % giving:

⎛⎞aaIB ETBE rmkETBE =−⎜⎟2 (2.37) ⎝⎠aKaEtOH ETBE EtOH

12 ⎛⎞−87.2 k =×1.209 10 exp⎜⎟ (2.38) ⎝⎠RT

Fite et al (1994) found that their data for ETBE formation were statistically and thermodynamically best represented by:

⎛⎞aETBE kaa⎜⎟IB EtOH − ⎝⎠K rETBE = 3 (2.39) aEtOH

12 ⎛⎞−10360 k =×4.7 10 exp⎜⎟ (2.40) ⎝⎠T

14580 lnKTT=− 1140 − 232.9ln() + 1.087 T (2.41) −×1.114 10−−03TT 2 + 5.538 × 10 07 3

Chapter 2: Literature Review 28 They obtained an apparent activation energy of 86.1 kJ/mol.

A comparison of these systems shows that Eley-Rideal type kinetics predominate. The reacting alcohols or water are adsorbed onto the active sites or centres of the catalyst and the olefins react with these from the liquid phase. Products did not appear in the denominators of these models, suggesting reaction and desorption as a single step.

The activation energies obtained for these reactions were quite high, falling in the range 80 to 95 kJ/mol. Given the use of Arrhenius type temperature dependence, this means that the reaction rates are very sensitive to temperature. As such reaction rate is s strong function of the catalytic distillation variable of column pressure.

2.3-5 Use of Solvents for Isobutylene Hydration

The use of a solvent for the purpose of accommodating the hydration of IB over Amberlyst15 may have several advantages according to the patent obtained by Moy and Rakow (1978): 1. improve the solubility of IB in a multiphase reactor, 2. improve miscibility of water and IB in a reactor operated under pressure, 3. push the equilibrium constant of the reaction towards TBA, 4. alter the sorption effects and diffusivity and reduce the extent of product inhibition. 5. provide a novel means of TBA separation.

Delion et al (1986) scrutinised the effect of a selection of common organic solvents upon experimental equilibrium constants. They also considered solvent inertness. The work showed that a ten-fold improvement in the equilibrium constant could be obtained by appropriate choice of solvent. The solvents, p-dioxane, acetone and nitromethane were considered inert. While alcohols such as isopropanol, acetic acid and the n-butyl mono ethylene glycol ether were seen to react, however it was noted that conditions could be kept such that this reaction was minimised while still allowing equilibrium to be measured. While the type of solvent used was seen to influence the extent of equilibrium stretching the investigators did not attempt to initiate the development of a model to predict this behaviour.

Chapter 2: Literature Review 29 A very useful study in terms of direct comparison to isobutylene hydration, is that conducted by Gonzalez and Fair (1997) on the hydration of isoamylene over Amberlyst 15 in the presence of acetone. They required the use of a solvent to keep the reacting mixture miscible especially under pressurised conditions of 600 kPa. They found that the solvent affected the activities of the reaction mixture in a favourable manner. As the acetone content increased the activity coefficient of water decreased and that of isoamylene increased. This boosted the formation of TAA as they found that water actually inhibited their reaction as mentioned previously.

The solvents identified as being of interest in terms of extractive distillation are those found in the mono glycol series such as ethylene glycol, propylene glycol, 2-methyl- 1,3-propanediol etc. The mono refers to one glycol as opposed to an ether type combination or polymer of glycol . Glycols or diols are compounds containing two alcohol groups. A US patent was obtained in 1978 for the hydration of isobutylene at high pressures (5 to 40 atm) over an ion exchange catalyst in the presence of solvents such as the glycol propylene glycol and various glycol ethers, Moy and Rakow (1978). They found that a minimum solvent to water wt ratio of 5 is required and that water to isobutylene mol ratio of 6 was optimum.

Glycols, however can react with isobutylene over solid acid catalyst such as Amberlyst 15 to form ethers. The complete kinetics of ethylene glycol or propylene glycol etherification with IB has yet to be published in open literature. The possible ethylene glycol ethers are formed by consecutive reactions as follows:

k1 HOCH2CH2OH + (CH3)2CCH2 ←⎯⎯⎯⎯→ HOCH2CH2OC(CH3)3 (2.42) k−1

k2 HOCH2CH2OC(CH3)3 + (CH3)2CCH2 ←⎯⎯⎯⎯→ (CH3)3COCH2CH2OC(CH3)3 (2.43) k−2 A preliminary investigation into the kinetics of the above reactions over Amberlyst 15 was conducted by Jayadeokar and Sharma (1993), in which they assumed elementary power law models. The rate constants they found are given in Table 2.08.

Chapter 2: Literature Review 30 Table 2.08 The kinetics of ethylene glycol etherification with isobutylene, at 60 oC, 1.5 atm, 10 wt % loading Amberlyst 15. Reaction k forward (s-1) k reverse (s-1) K

Rxn 1. Eqn ZZZ 1.437 × 10-5 6.361 × 10-5 0.2259

Rxn 2. Eqn ZZZ 2.848 × 10-6 1.891 × 10-5 0.1506

Knifton (1994) outlined the use of glycols as an alternative to MTBE for capturing IB. He found that propylene glycol was more selective to the mono ether than ethylene glycol. He also proposed that some glycol ethers might find use as fuel oxygenates.

2.4 Extractive Distillation Technology

In equivalence to reactive systems the addition of solvents to separation systems such as distillation and the subsequent modification of activities can also be beneficial and enhance performance. Extractive distillation is defined as distillation in the presence of a miscible, high boiling, solvent that does not form any new azeotropes in the system and alters component relative volatilities to such an extent as to allow azeotropes to be ruptured and close boiling mixtures to be separated. In extractive distillation it is the component having the greater volatility, not necessarily boiling point, which is rich in the distillate. Thus the natural tendency for separation in distillation can be reversed. The other component leaves in the bottoms with the solvent and these are passed to a second column were the solvent is recovered, Figure 2.01. Solvent promoting natural separation tends to result in a more economic process, as less of it is required.

The solvent modifies the relative volatility α12 by affecting the ratio of the liquid phase activities. The relative volatility at low to moderate pressure can be defined as:

vap yxik γ iiP a12 ==vap (2.44) xyik γ kkP

vap At higher pressures the Pi term is replaced with standard state fugacity. In terms of obeying the general rule of promoting natural tendency, a solvent that increases γi vap vap relative to γk, when Pi > Pk , is favoured over one that increases γk relative to γi. To force the naturally more volatile component i overhead, the solvent should either:

Chapter 2: Literature Review 31 Case 1. behave essentially ideally with component k and cause positive deviation from

Raoult’s law for component i (γk ≈1 and γi >1). Case2. behave essentially ideally with component i and cause negative deviation from

Raoult’s Law for component k (γi ≈1 and γk <1).

Compounds of similar shape and size, eg. pentane-hexane, tend to behave ideally with γ ≈1. Dissimilar compounds repel each other with positive deviation and compounds that tend to associate, due to polarity and hydrogen bonding in the liquid phase, cause negative deviation. As systems exhibiting positive deviation are more common, the usual approach is to force the lower boiling component overhead by selecting a solvent that is chemically similar to the higher boiling species and dissimilar to the lower boiling one.

The distillation of ethanol or TBA with water in the presence of ethylene glycol is an example of Case 2., where water associates with ethylene glycol. Thus water will experience negative deviation, γk <1, and TBA will interact ideally with ethylene glycol and its activity coefficient will approach 1. The implication of this in terms of reaction and catalysis are yet to be mapped out. The effect upon equilibrium is likely to be that of pushing it to the right towards TBA formation.

The process of extractive distillation has been defined as a combination of absorption and distillation by Stichlamir and Herguijuela (1992). The first column can be identified as an absorption column as the low boiling azeotropic fraction is fed near the bottom and the high boiling entrainer is fed pure at the top, and additionally both flowrates are higher than that of the reflux or reboil. Further the solvents can be chosen according to the rules of absorption. Such a definition of the process can lead to simplifications in testing and experimental arrangement. For instance to allow for ease of laboratory experimentation and to be able to work in a near continuous mode as opposed to batch extractive distillation. Resa et al (1995) devised specialist equipment making use of this representation. A bypass was installed above the reboiler such that the bottom liquid flow was taken off as the bottoms and not allowed to return to the reboiler. The reboiler itself was charged with azeotrope. Using this arrangement they were able to determine extractive distillation VLE of acetone-isopropyl ether, Resa et al (1995), as well as carry

Chapter 2: Literature Review 32 out an entrainer screening program for this system, Resa et al (2000). The equipment devised is essentially an absorption column and greatly simplifies the experimental column setup.

2.4-1 Entrainer Selection

The selection of a good solvent is the most important step in developing a successful and viable extractive distillation process. The requirements of a satisfactory extractive distillation solvent are: 1. Ability to alter the VLE of the original mixture with only small quantities of solvent. 2. Absence of azeotropes between the solvent and components of the mixture. 3. High capacity, or ability to dissolve the components of the mixture. 4. Low volatility to prevent vaporisation with the overhead product and to maintain high concentration in the liquid phase. 5. Separability, the solvent must be readily separated from the mixture to which has been added.

Various methods are available for screening of potential solvents and estimation of their effectiveness. These include residue curve maps, relative volatility at infinite dilution predicted and determined experimentally and solvent free x-y diagrams.

Residue curves are particularly useful if thermodynamic parameters are available for the compounds of interest and provide an excellent tool for initial rapid scanning especially if a simulation package such as Aspen Plus is available. The generation of residue curves and their interpretation for advanced distillation methods was described in detail in the review on azeotropic distillation given by Widagdo and Sider (1996). Using residue curves the occurrence of azeotropes can be quickly discounted and effectiveness evaluated through simple graphical means as developed by Bauer and Stichlmair (1995). A line is drawn through the inflection points of the residue curves from the binary azeotrope to the binary edge of the entrainer and component, which appears overhead, Figure 2.01. The better the solvent, the further away from its corner does the line intersect the edge. Thus entrainers can be quickly evaluated in terms of effectiveness and ability to promote natural separation.

Chapter 2: Literature Review 33

Figure 2.01 Residue curves for the extractive distillation of aceton and methanol using either water or ethylbenzene, Bauer and Stichlmair (1995).

2.4-2 The TBA and Water Azeotrope

The azeotrope occurring between TBA and water is pressure insensitive and its location in either composition or temperature does not shift significantly with variation in pressure, see Table 2.09 Horsley (1984). The implication of this is that the azeotrope can not be broken by merely adjusting the column pressure in a conventional distillation system.

Table 2.09 Insensitivity of water/TBA azeotrope towards pressure Pressure (atm) TBA bp oC Azeotrope bp TBA mol % Az oC 0.132 82.5 79.9 64.57 10.684 82.9 79.85 59.87

The central location of the azeotrope makes it stable and difficult to shift. The azeotropic behaviour of TBA and water in general are quite similar. There are few compounds, which form azeotropes with one and not the other.

2.4-3 Extractive Distillation Applications

The use of ethylene glycol and diethylene glycol for the dewatering of a range of alcohols in the C1 to C4 range was investigated by Liu et al (1993). Their study is encouraging as they found good selectivity in terms of ethylene glycol rupturing the

Chapter 2: Literature Review 34 TBA/water azeotrope. The values of selectivity for TBA/water system were actually higher than those for the ethanol system, see Table 2.10. If accurate this would allow for lower usage of glycol and improved economics. Keeping the solvent fraction low may also be important in preventing side reaction.

Table 2.10 The selectivity of ethylene glycol for EtOH and TBA

System Xent=0.0 Xent=0.5 Xent=0.7 Xent=0.9 Value: a12 (a12)ent/a12 (a12)ent/a12 (a12)ent/a12 EtOH-water 2.79 3.5 2.89 2.02 TBA-water 2.19 4.01 5.71 5.96

Berg and Yang (1992) applied for a US patent focusing on TBA dewatering by extractive distillation. They proposed an extensive list of agents, which included various benzoates and glycols. Some of the proposed solvents violate the requirements of extractive distillation solvent and form azeotropes. They ran examples in both an equilibrium still and a small pilot scale column and evaluated relative volatilities in order to grade their solvents. The relative volatilities they calculated are given in Table 2.11, for the solvents that gave TBA as overhead of the first column.

Table 2.11 Potential Entrainers for the separation of water and TBA, Berg and Yang (1992).

Solvent a12 triethylene glycol 2.33 polyethylene glycol 400 1.92 dimethyl sulfoxide 1.51 dimethyl phalate 1.41 1,4-butanediol 1.38 dimethyl formamide 1.35 dipropylene glycol 1.32 2-methyl-1,3-propanediol 1.32 tetraethylene glycol 1.28 diethylene glycol 1.26 methyl benzoate 1.21 dimethyl adipate 1.21 1,3-butanediol 1.2

They were obviously unable to include common extractive distillation solvents such ethylene and propylene glycol. The di- and tri- glycols are actually ethers with a central oxygen, eg. diethylene glycol has the formula: O(CH2CH2OH)2. The authors provided a table of the likely column size required, Table 2.12.

Chapter 2: Literature Review 35

Table 2.12 Predicted column dimension given a12, Berg and Yang (1992) * ** a12 N theoretical N actual 1.2 52 70 1.5 23 31 2.0 13 17 2.5 10 13 *theoretical equilibrium plates required at total reflux to gain 99 mol %. **actual plates required at 75 % overall plate efficiency.

The most comprehensive published pilot plant study on ethanol dewatering was conducted in Brazil by Meirelles et al (1992). Where ethanol is produced from sugar cane biomass waste and is blended in very high ratios with petrol and thus requires high purity and dryness to prevent mixture separation and poor engine performance. A laboratory column was first used with a diameter of 65 mm and 24 bubble cap plates. Pilot plant columns of 80 mm diameter were subsequently employed. The extractive column contained 60 sieve plates with downcomers and feed located on plate 31 from top. The solvent recovery column contained 40 plates and feed located on plate 19. The columns achieved top compositions of more than 99.5 mol % ethanol. The temperature profile was flat, with one zone between the entrainer feed location and the column feed and the other below the column feed.

They observed that solvent to feed ratio was the most significant variable and performance for a particular S/F rate change little with substantial reflux manipulation, 0.71 to 2.01. They recommended a moderate S/F ratio such as 0.71 and a minimum practical reflux of 0.5, to be the most economic for feeds of 15 mol % water in ethanol. The reflux of 0.5 was found to be the minium practical reflux beyond which performance dropped off.

They found that solvent water content could substantially degrade the purity of the ethanol, which follows from the fact that the solvent is fed near the top of the column. In terms of operation, extractive distillation proves to be simpler than azeotropic distillation as the columns used are largely independent of each other. As with reactive distillation the columns can be controlled based on column temperature profile.

Chapter 2: Literature Review 36 Recent advances in residue curve map technology and the interest in breaking azeotropes occurring in catalytic distillation are bringing the two fields closer together. Given its relative simplicity in application for overcoming azeotropes and established use in industries such as alcohol production, extractive distillation entrainers should find use in reactive and catalytic distillation systems.

2.5 Catalytic Distillation Technology

Catalytic distillation is a new and exciting concept, which has attracted the attention of many reactor engineers. It is an example of unit intensification where reaction and separation are combined with the aim of improved efficiency and better plant economics. The potential to push beyond equilibrium yields, heat integration and capital cost reduction are all very desirable. The field of reactive distillation has grown significantly in the last 30 years. A total of 562 publications for the period 1970-1999 are listed in the Engineering Index and a total of 571 US patents for the period 1971- 2000, Malone and Doherty (2000).

When the temperatures of reaction and separation overlap, and there is good thermodynamic reason, catalytic distillation becomes an attempt to juggle the requirements of separation and reaction through hardware design and operating variables. Although the number of applications keeps increasing as demonstrated by Hiwale et al. (2004) in their review of industrial applications. It has been recognised that not all systems are suitable, feasibility studies are very important and hence the growing interest in such tools as reactive residue curves Thiel, Sundmacher and Hoffmann (1997). Hindrances to full development can arise due to side reactions as found recently with attempts in applying MTBE to isobutylene capture and purification in coupled column process, Qi et al (2002). The presence of reactive and conventional azeotropes can limit the success of unit intensification as found for butylacetate production via catalytic distillation where the successful design ultimately results in a train of columns, Hanika et al (1999). The occurrence of multiple steady states and the susceptibility of some processes to these can lead to problems with startup and controllability. Further, contradicting design requirements can be found between the amount of catalyst required to prevent frequent repacking due to catalyst deactivation and slow reaction rate and the requirement of good distribution to allow for good mass transfer and separation

Chapter 2: Literature Review 37 efficiency Krishna (2002). An understanding of possible challenges and limitations is growing in a field, which is rapidly maturing.

A catalytic distillation column is often divided into three distinct sections rectifying, reactive and stripping. The proportions and feed location change depending on reaction and separation demands. The packing and width of the three sections is often different, Subawalla et al (1997). The choice of internals and their design can greatly influence feasibility, yield and purity. As these depend on reaction rate, separation mass transfer rates the liquid holdup and catalyst concentration. Hardware design and good characterisation is increasingly coming of importance especially as new applications are investigated Gotze et al (2001) and modelling becomes more representative and complex Taylor and Krishna (2000).

The major fields of research in catalytic distillation can be categorised as: • New applications and process improvement (cf Section 2.5-1) • Process synthesis and design tools (cf Section 2.5-2) • Modelling and simulation (cf Section 2.5-3) • Hardware design (cf Section 2.5-4) • Dynamics and control (cf Section 2.5-5)

Each of these fields, which are continuously being advanced, need to be utilised to obtain a successful catalytic distillation design. Tuchlenski et al (2001) demonstrated how industrial applications draw upon each developing field in their evaluation and application of an industrial column. This section of the literature review deals with each of these fields in turn, with initial process design and feasibility in mind.

2.5-1 New Applications and Process Improvement

Hiwale et al (2004) presented a near exhaustive listing of the various applications of catalytic and reactive distillation up to the date of their review. Applications have expanded beyond esterification and etherification to hydrolysis, , transesterification, alkylation and various other organic reactions. The development of new applications has particularly gained momentum in the last five years.

Chapter 2: Literature Review 38 The expansion to new applications has been so rapid and widespread as carrying out reaction and separation in one piece of equipment can offer several advantages over a conventional reactor and distillation train. The potential advantages include: 1. Unit intensification and the reduction in plant cost: if separation systems can be reduced and simplified then this can lead to significant capital savings, Tuchlenski et al (2001). 2. Shifting of equilibrium: by continuously removing the products as the reaction proceeds. Equilibrium limited reactions can be made to give higher conversion than can be attained in a conventional reactor. 3. Heat integration: the heat generated by exothermic reaction is used to assist the distillation duty. The heat is transferred directly at very high efficiency. 4. Rupturing of some azeotropes: some azeotropes can be reacted away some can be avoided by feed location and the action of reactants as entrainers such as for the reactive distillation of methyl acetate, Towler and Frey (2000). 5. Improved selectivity: if a reactant or product involved in subsequent reactions can be removed from the reaction zone, such as in the adol condensation of acetone to acetone alcohol where the series formation of mesityl oxide can be avoided Podrebarac et al (1998).

The potential disadvantages: 1. Residence time requirements: if the residence time required for the reaction is long the column may be very large but produce very little. 2. Plant scale: it is difficult to design a reactive distillation column for large flowrates. 3. Process incompatibility: due to volatility constraints or temperature requirements of the reaction being different that of the separation. 4. Control Aspects: unit intensification reduces the degrees of freedom and it may be difficult to achieve the desired yield and purity due to control constraints, Sneesby et al (1999). 5. Mass transfer limitations using current hardware designs can limit the benefits of catalytic distillation as in the case of adol condensation of acetone, which enjoys better product selectivity at the expense of sizing and catalyst usage, Nicol (2003).

Chapter 2: Literature Review 39 Given that the interest here is the application of CD for the synthesis of TBA particular applications have been focused on, which may give insight to the problem at hand. By understanding what makes these similar applications successful and where their limitations lie more informed design decisions should be able to be made in regrades to design of catalytic distillation for synthesis of TBA. The same systems considered for similar reaction kinetics will be considered here. The systems which were considered similar include those for the synthesis of: MTBE, ETBE, TAME, TAA and TBA dehydration.

2.5-1-1 Methyl tert-Butyl Ether, MTBE

It is fortunate that the MTBE system can be considered as being related to that of TBA as it is the most researched system in reactive distillation. The banning of MTBE as a fuel additive in the United States of America has not stemmed the flow of papers exploring the process or using it as an example to develop advanced topics in catalytic distillation. Some of the topics explored have been those of side reaction, multiple steady states, reactive azeotropes and various aspects of catalytic distillation modelling including the development of rate based models.

In order to determine what makes the MTBE system successful early experimentation in kinetics and column operation was refereed to. A large part of the success of the MTBE system can be attributed to the fast kinetics within a miscible system over acidic resins. The most widely referred to and the most comprehensive experimental study in MTBE catalytic distillation is that of Flato and Hoffmann (1992). They presented a detailed account of design, procedure and range of product attainable for a complete catalytic distillation column for the synthesis of MTBE. The MTBE process as described clearly demonstrated the potential advantages listed above. They compared the reactive distillation process against established industrial processes (Hüls AG, Snampogetti SpA, IFP, Texaco AG, Ec-Edölchemie) for the synthesis of MTBE and showed a clear reduction in unit operations required. They demonstrated equilibrium shifting and heat integration.

A key factor to successful operation of the system is adequate pressurisation. Relatively mild pressures are required to liquefy isobutylene, 2.5 bar at 25 oC, however more

Chapter 2: Literature Review 40 importantly at pressures in excess of 4.03 bar the volatility of MeOH and MTBE are reversed and MTBE becomes the heaviest component in the system. The higher the pressure the wider the gap between the two and the greater the separation potential. Forcing MTBE to be the bottom product provides leverage for dealing with the azeotropes of the system. Commonly encountered azeotropes are minimum boiling. If the component of interest is not the major component of such an azeotrope and is the heaviest boiler of the system than it should be able to be withdrawn at high purity as the bottoms. This principle has also been employed for ETBE, TAME and TAA catalytic distillation design.

In their experimental investigation, Flato and Hoffmann (1992), varied the isobutylene feed mass fraction as 15, 30 and 50 wt % of a mixed isobutylene/butenene-1 feed in representation of a typical C4 fraction. They employed Raschig rings manufactured of ion exchange resin material in a pilot scale column for the synthesis of MTBE. The catalyst Raschig rings of 6 mm diameter showed an ion exchange capacity of 4.52 mol H+.kg-1 which compares well to that of Amberlyst 15 of 4.9 mol H+.kg-1. The column employed was of 53 mm internal diameter and 1600 mm long. The ion exchange Rachig rings where placed in the top half of the column (5100 mm bed) and non-catalytic rings were placed in the bottom half (5100 mm bed). The column itself was of a special flangeless design to minimise local heat loss.

Best results were obtained with a MeOH/IB feed ratio of 0.8 to 1.2, especially in terms of selectivity towards MTBE relative to the side reaction of isobutylene to diisobutylene. Feeding MeOH above the reactive zone and isobutylene below allowed conversions in excess of equilibrium to be gained with IB 30 wt %, MeOH/IB of 1 and pressures of 8 and 9 bar. A mixed feed above the reactive zone could not achieve conversion in excess of equilibrium. Higher pressures gave nearly complete conversion of MeOH to MTBE as a result of higher reaction rates and improved fractionation potential.

Sundmacher and Hoffmann (1994) and (1996) devoted considerable effort to improving the quality of the catalytic Raschig ring packing used and gaining absolute understanding of the system. They probed the catalytic distillation process employing

Chapter 2: Literature Review 41 experimental design and examination of experimental results through modelling. Sustained oscillations were observed in temperature, pressure and in reflux flowrate for both non-reactive and reactive distillation. These were attributed to the distillation behaviour of the binary mixture IB/MeOH, and could be avoided by adjusting the feed flowrates. They found that the catalytic Raschig packing produces MTBE under trickle flow conditions. That the resistance of the liquid film surrounding the catalyst rings is negligible and that the main mass transfer resistance at the vapour/liquid interface is located on the liquid side. They incorporated microkinetic and macrokinetic rates initially for an EQ model and then for a detailed NEQ model. The modelling was conducted using Maxwell-Stefan mass transfer equations.

Once the kinetics had been defined and the process demonstrated experimentally the catalytic distillation of MTBE was extensively modelled with increasingly higher accuracy and complexity of models. Nijhuis, Kerkhof and Mak (1993) identified the existence of multiple steady states (MSS) within the MTBE system. MSS in a reactive distillation column refers to output multiplicities. That is, a column of a given design exhibits different column profiles at steady state for the same set of inputs and the same values of operating parameters.

Nijhuis, Kerkhof and Mak (1993) found MSS by varying the feed location of methanol to an MTBE catalytic distillation column. Thus obtaining multiplicity through design parameter manipulation. They found that increasing catalyst loading towards equilibrium conversions had only a minor effect upon MSS. Decreasing catalyst loading eventually resulted in a single steady state. Thus in their opinion the MSS occurred when operating close to equilibrium and were considered unlikely to be a result of kinetic limitation. The authors pointed towards the exothermicity of the chemical equilibrium reaction to be responsible for MSS.

Jacobs and Krishna (1993) also found multiple solutions by varying the feed location of methanol to a MTBE column modelled in AspenPlus. They were unable to tie the observed effects to the crossing of non-reactive distillation boundaries. They found that manipulating catalyst loading towards reaction equilibrium did not effect the observed MSS. Thus they discounted kinetic based multiplicity as had been found possible for the

Chapter 2: Literature Review 42 system at hand by Rehfinger and Hoffmann (1990). What they did notice upon constructing reactive distillation residue curves of the system was that the starting point of a residue curve influenced the MSS attained. High conversion corresponded to residue curves starting near the n-butene-methanol azeotrope, low conversion corresponding to residue curves starting in pure isobutylene.

Hauan, Hetzberg and Lien (1995) clearly demonstrated the ability of n-butene to influence low and high steady state. The dilution effect of n-butene was seen as beneficial as it lowers the reaction zone temperature improving the equilibrium conversion. Further the effect of n-butene upon the activity coefficients is positive in respect to the kinetic model employed. Two conditions were identified that were seen to influence whether high or low steady states were attained. Firstly the reaction mixture in the lower part of the reactive zone must be sufficiently diluted so that MTBE is not decomposed before it escapes the reaction zone. This highlighted the classic concern of carrying out a reaction of the type:

cat A + B ←⎯⎯⎯⎯→ C (2.45) within a catalytic distillation environment. That separation within the reaction zone would simply lead to product loss through reverse reaction. Secondly when methanol is fed below the reactive zone sufficient amounts of other components including MTBE must be present to lift the methanol up to the reaction zone. The internal recycle of MTBE mechanism proposed was strengthened through a dynamics study conducted by the Hauan et al. (1997). They showed that small changes typically resulted in sustained oscillations, resulting from the recirculation effect and that further larger changes could bring about a drop to the lower conversion branch.

Mohl et al (1999) presented the first rigorous experimental verification for the existence of multiple steady states in catalytic distillation for MTBE and TAME and developed a strong argument in support of kinetic instability as cause of MSS. They investigated the operating parameters reflux ratio R and reboiler duty Q (kW), which they considered to be different to investigation of design parameters. They demonstrated that for the MTBE system MSS occur within a very narrow region defined by R and Q. For MTBE generally the region commenced at R = 10. For R = 20 the widest point of the wedge

Chapter 2: Literature Review 43 defined gave a range of 0.1 kW. Thus experimentally it was found difficult to verify multiplicity for MTBE. The authors also studied the TAME system and found it more susceptible to MSS and thus MSS more attainable experimentally. The region of multiplicity was broad and the experimental conditions could be reproduced with sufficient accuracy to allow for experimental validation. They achieved two different steady state temperature profiles by varying startup procedure. The low steady state was obtained by charging the reboiler with the azeotropic mixture of methanol and n- pentane (the inert). This reinforced the residue curve findings of Jacobs and Krishna (1993). They also demonstrated experimentally that a pulse disturbance of TAME could create a transition between steady states. The authors drew the conclusion that for catalytic distillation overall it may be difficult to reach the desired steady state at startup and that a temporary disturbance could cause a transition to an undesirable steady state.

Mohl et al (1999) give a physical explanation for the steady states they observed. They began by identifying that for standard reactors multiple steady states arise due to: 1. heat effects (adiabatic operation) 2. kinetic instabilities

Further that for distillation columns multiple steady states can occur as a result of: 1. dependence of the molar volume or the heat of vaporisation on the composition for special choices of input variables 2. special topology of the residue curves for azeotropic mixtures.

Thus for reactive distillation it is possible that multiple steady states can arise due to any of these, as well as their combination and further by some mechanism arising from the interaction of reaction and separation. They narrowed their search to kinetic instability and in particular the self-inhibiting aspect of the TAME kinetic expression. They showed that as the number of active sites is increased and the process gets closer to reaction equilibrium the range of heating rates for which multiple solutions are obtained decreases and finally vanishes. Thus moving away from the kinetically controlled region to the equilibrium controlled column, moves away from the possibility of multiplicity for operating parameters. Using a CSTR and a single staged reactive distillation column, with the same kinetic model they demonstrated that an overlap of

Chapter 2: Literature Review 44 multiple solutions regions exists in the Damköhler number ((Da), cf Section 2.5-2-1) and initial molar ratio. They found that if the same number of active sites were modelled with more stages and greater separation the region rapidly expanded.

Baur, Taylor and Krishna (2001) developed a NEQ dynamic model and showed that transitions between the steady states for MTBE and TAME were influenced by the internal isobutylene recycle flows established upon a feed flow perturbation. Developing their model further they reapplied it to the data of Mohl et al (1999) and were able to capture the same transitions. They found that the transient behaviour is very sensitive to the static liquid holdup in the catalytic section. This reinforces the relation to Da and further stresses the importance of understanding of column hydrodynamics.

The product inhibition mentioned is common to the all the systems referred to in Section 2.3-4, MTBE, ETBE, TAME, TAA and TBA. Thus some degree of multiplicity may be expected in each. It is possible to avoid multiplicity for these systems through manipulation of Da. The explanation given by Mohl et al however is not exhaustive as Jacobs and Krishna (1993) found that increasing the amount of catalyst did not effect multiple steady states. Further the MSS they observed occurred under conditions of reaction equilibrium. The study presented by Mohl et al largely neglects side reaction, which may also contribute to MSS. As has been mentioned a number of mechanisms have been suggested for MSS including heat effects such as that of exothermic reaction, kinetic instability and the crossing or existence of distillation boundaries. Further considerations have been raised such as the interaction of non-reactive sections with reactive sections and the strong interaction of the two feeds methanol and mixed C4, Guttinger et al (1999). It is likely that in the highly non-linear system multiple causes can give multiple types of steady states, reached through design or operating parameters.

Endeavours to gain a better understanding of CD for the synthesis of MTBE saw the rapid development of modelling, the desire to extend catalytic distillation technology to different systems resulted in the development of reactive distillation thermodynamics and feasibility determination. In particular the development of residue curve theory was

Chapter 2: Literature Review 45 based on the knowledge gather for the MTBE system. Venimadhavan, Buzard, Doherty and Malone (1994) established the reactive distillation residue curve and demonstrated its use for the homogeneously catalysed MTBE system. The method developed was customised for heterogeneous catalysis and the residue curves for MTBE and TAME drawn by Thiel, Sundmacher and Hoffmann (1997). It was realised that through manipulation of pressure, holdup and vapour rate the time available for reaction relative to separation could be altered such that the residue curve moved away from one defined by separation features such as distillation boundaries to one defined predominantly by reaction equilibrium. This affect can be clearly seen in Figure 2.02 as the residue curves converge to the equilibrium curve with increased Da. The conventional azeotropes shown disappear and new stationary points appear and were labelled kinetic pinches. Although the azeotropes are no longer represented it is difficult to establish the physical interpretation of their absence, especially given that others have used their presence to explained the observed experimental and modelling effects Hauan et al (1995). In residue curve analysis while kinetics and equilibrium are incorporated, mass transfer effects are notably absent. Mass transfer over the packing and within the particles of catalyst is not included as part of the definition of the residue curves created.

Chapter 2: Literature Review 46

Figure 2.02 Reactive residue curve for heterogeneously catalysed MTBE synthesis at operating pressure P = 0.8 MPa; Damköhler number (a) Da = 0 (b) Da = 10-04 (c) Da = 2 × 10-04 and (d) Da = 1, Thiel et al (1997)

One advantage of carrying out the MTBE reaction in a catalytic distillation column is that the isobutylene can be fed along with other components of the C4 cracking stream and these leave the top unreacted. Lately this advantage has been further utilised in processes, which treat methanol, as a reactive entrainer forming MTBE only with isobutylene allowing the remaining C4 components to be separated. The MTBE is decomposed in a second column to give pure isobutylene for use downstream in say alkylation or polybutylenes production, Qi et al (2002). The side reactions possible forming water, diisobutylene and dimethyl ether do detract from the successfulness of two coupled columns and result in the methanol recycle having to be purified. Even if side reactions are kept to a minimum the enrichment of their products through recycle

Chapter 2: Literature Review 47 will be an important consideration for any process such as catalytic extractive distillation employing a solvent.

The qualitative consideration of the decomposition column for such a process given by Beckmann et al (2002) gives good industrial insight into the application of catalytic distillation technology (authors being representatives of Oxeno Olefinchemie, Degussa and Sulzer Chemtech LTD). From industrial experience they suggest that catalyst deactivation for MTBE processes is low, with a typical batch of ion exchange resin giving 8000 h of high stability. They found that the EQ model used gave a good fit to laboratory data but failed for the pilot plant process. They suggest that Katapak-S at a laboratory scale does not correspond very well to Katapak-S at the technical scale. They recommend pilot plant evaluation on the basis that NEQ models, which may give a better fit, require the evaluation of mass transfer coefficients and dispersion coefficients for which the attainment of required accuracy is an ambitious task.

Although Flato and Hoffmann (1992) demonstrated the use of a single column to attain MTBE it has since been found that the use of a pre reactor may give the most economically viable process solution. Both the CDTech process and the Ethermax process by UOP consist of fixed bed reactors followed by a catalytic distillation column to reach high conversion. The use of a prereactor should be considered when the approach to equilibrium in terms of reaction rate is steep and only dies out right before equilibrium is reached. The behaviour of the MTBE formation rate falls in this category and a prereactor is generally used, approaching equilibrium to within 5 % in IB conversion. The prereactor conversion is typically 88-92 % and the conversion of the reactor and column approach 100 %, Subawalla et al (1999).

2.5-1-2 Ethyl tert-Butyl Ether, ETBE

The most significant advantage of ETBE over MTBE is the renewability of ethanol produced by fermentation of biomass. ETBE has a higher octane rating and lower volatility as well as being less hydrophobic. However its cost of production and the amount required for blending are higher. Reactive residue curves were constructed by Thiel, Sundmacher and Hoffmann (1997) in much the same manner and with the same qualitative results as those developed previously for MTBE and TAME.

Chapter 2: Literature Review 48

Sneesby et al (1997) & (1997) have modelled the process with particular attention paid to control and dynamics. Ethanol and ETBE enjoy the crossing over of vapour pressure lines in much the same manner as methanol and MTBE. The pressure required to gain cross over is 2.2 atm at a shared boiling point of 100 oC. Amberlyst 15 has an operating temperature limit of 120 oC. ETBE boils at 120 oC at 3.63 atm and EtOH boils at 120 oC at 4.24 atm. Through the dilution effect of isobutylene and n-butene Sneesby et al achieved the distribution 75 to 81 oC across the reaction zone operating at 9.5 atm. The reboiler temperature was 159 oC and the condenser temperature 74 oC. As in the case of MTBE higher pressures give a better separation potential in the bottom of the column. Conversions of 98.3 mol % were achieved in isobutylene. A purity of 96.1 wt % ETBE was gained in the bottoms with EtOH the main contaminant. Sensitivity analysis found the conversion and purity to be very sensitive to reboiler duty. Multiple steady states were not encountered but their likelihood was mentioned. Controlling for both purity and conversion was seen as creating a control conflict.

2.5-1-3 tert-Amyl Methyl Ether, TAME

The TAME system has been demonstrated to have a high potential for multiple steady states as demonstrated by Mohl et al (1999). In their pilot plant investigation of the application of catalytic distillation to TAME synthesis Bravo, Pyhalahti and Jarvelin (1993) pointed out operational difficulties. Multiple steady states were experienced experimentally, with startup conditions determining the state attained. Their activity coefficient calculations showed that in this system MeOH had an activity coefficient of 4 to 6 while the other components had coefficients close to unity. Thus MeOH was highly non-ideal with respect to the other components. Different methanol to olefine feed ratios produced different operating regimes. The difference in MeOH and C5 olefine volatilities made it difficult to maintain optimum concentration within the reactive zone. They also point out that they encountered a wide range of side reaction products including dimethyl ether and tert-amyl alcohol. Beyond a consideration of activities they found that flat concentration profiles gave experimental indication of MSS. The authors struggle to justify reactive distillation for the synthesis of TAME.

Chapter 2: Literature Review 49 Baur and Krishna (2002) and Subawalla and Fair (1999) both chose this system to develop design guide lines for reactive distillation. Baur and Krishna compared the designs attained when catalytic Raschig rings or Bale packing were employed. The design method of Subawalla and Fair is high methodical and they do obtain a successful design through modelling. They recommend the use of a prereactor on the basis that the reaction is slow and requires much methanol and catalyst.

2.5-1-4 tert-Amyl Alcohol, TAA

Interest has been expressed in developing a catalytic distillation process for tert-Amyl Alcohol production by direct hydration of isoamylene over heterogeneous catalyst such as Amberlyst 15. TAA is the equally troublesome sister of TBA. Smith (1980) developed Bale packing, and patented his C4 purification process making use of it. A decade later he patented his own research into olefine hydration for the production of TBA and TAA, Smith (1991). Neither alcohol application has yet to be developed and only the presence of TBA and TAA was confirmed. The researchers group of J. R. Fair at the University of Texas Austin chose to investigate the synthesis of TAA over Bale packing Gonzales et al (1997). The product TAA having the highest boiling point allows a column operating in near total reflux with TAA being drawn off the bottom.

Isoamylene and water are largely immiscible and need to be brought into contact over the chosen catalyst in order to gain effective rates of reaction. Further, the limiting reagent isoamylene is the most volatile component and its volatility is substantially different from that of water. In light of this Gonzales et al (1997) proposed the use of acetone as solvent for the reaction. The column was operated in total reflux to minimise reactant loss. A pressure of 250 kPa was employed and TAA was withdrawn at the bottom. The reactive zone was located at the top of the column given that the limiting reagent isoamylene is the most volatile component. Further the isoamylene was fed at the bottom of the column and the water at the top.

González, Subawalla and Fair (1997) observed some side reaction of acetone to mesityl oxide, however they claimed very low yields especially in comparison to TAA. However the experimental evaluation was conducted with average values of 90 mol % acetone, 7 mol % water and 3 mol % isoamylene, in the feed. Thus the process was

Chapter 2: Literature Review 50 acetone rich and the extent of side reaction at 0.3 mol % yield would have consumed considerable amounts of the solvent. The solvent here needs to be vaporised and kept in the column, which creates an additional duty demand.

The azeotropic behaviour of this system is complex and the authors list the order of volatility as: 1. acetone-isoamylene azeotrope, 25 mol %, bp 32oC 2. isoamylene, bp 37.5 3. acetone, bp 56.5 oC 4. water-TAA azeotrope, 25 mol % water, bp 87.4 oC 5. water, bp 100 6. TAA, bp 102 oC

The azeotrope between water and TAA, which is rich in TAA, should have kept much of the TAA in the column once it was formed. Yet a good proportion of TAA ended up in the bottom of the column in both experiments and simulation. The fact that acetone alters the activities of the reacting mixture has been noted by the authors in their reaction kinetics study, Section 2.3-4. In light of this it can be suggested that the action of acetone in the system, beyond improving miscibility, is likely to be that of extractive distillation entrainer. A better choice of solvent and optimisation in terms of extractive distillation could allow for overall process improvement..

2.5-1-5 Dehydration of TBA

The dehydration of TBA is an obvious choice for catalytic distillation as isobutylene is instantly released to the vapour phase. TBA is maintained in the reaction zone and pure water withdrawn as bottoms. The dehydration of TBA is used to obtain isobutylene for the production of MTBE, isooctane and polyisobutylene. The advantages of carrying out TBA dehydration in a catalytic distillation column are mild conditions ~120 oC, quantitative TBA conversion per pass and the ability to use feedstocks of a lower quality, and most significantly at an industrial scale the prevention of catalytst poisoning by metal ions present in the feed.

Chapter 2: Literature Review 51 Knifton et al (2001) conducted experimental studies on the dehydration of TBA. Generally near pure isobutylene product and pure water are obtained, with sufficiently low TBA feed rates. Abella et al (1999) used a pure TBA feed at the bottom of the reaction zone. In a 25 mm column they used 120 mm of mesh saddles as a rectifying section, 220 mm of their Amberlyst 15 based pellets and 160 mm for a stripping section. In their simulation study this equated to two rectifying, five reactive and thee stripping theoretical equilibrium stages. They employed the UNIQUAC thermodynamic model and their own kinetics, Abella et al (1999).

It can be seen that for the synthesis processes mentioned above various common aspects can be identified. The order of volatility is important and generally the product is withdrawn at the bottom of the column. This is significant as one of the reactants is relatively low boiling. The order for MTBE and ETBE synthesis is forced through pressure manipulation. This also allows for dealing with azeotropes, which tend to be minimum boiling, which persists irrespective of reaction.

The isobutylene miscibility of the successful systems MTBE and ETBE is relatively high. The inert diluent 1-butene and 2-butene are significant. They alter activities and provide dilution. The use of solvents such as acetone can allow for solubilities approaching those in the miscible systems to be approached, however solvents can create reaction networks which need to be optimised for product selectivity.

2.5-2 Process Synthesis and Design Tools

Here the tools of feasibility evaluation as well as new case studies employing the tools being developed are considered. Quite obviously not all reaction systems are suitable for catalytic distillation and further the extent of advantages gained with some may not warrant its use. Deciding upon the suitability of a system and evaluating the feasibility of a catalytic distillation column design has become a major area of research and interest. The most powerful feasibility tools developed thus far are those of the field of catalytic distillation thermodynamics. These include reactive azeotrope determination techniques and reactive distillation residue curves.

Chapter 2: Literature Review 52 At the simplest level reactive distillation can be considered for processes where the reaction occurs in the liquid phase. Where the conversion is limited by unfavourable chemical reaction equilibrium, the temperature range of the catalyst activity overlaps well with the temperature profile of the separation and when the full advantage can be gained of combing separation and reaction in one unit, Westerp (1992).

An effective way of decomposing the design and development of reactive distillation involves four stages as proposed by Malone and Doherty (2000): 1. Feasibility and alternatives: the use of residue curves, reactive distillation parameters, literature review, possibly experimentation to complement literature review, similar processes. 2. Conceptual design and evaluation: generally involves short cut calculation methods, followed by simulation of gradually increasing complexity. 3. Equipment selection and hardware design: this is commenced when increasing simulation complexity, involves evaluating technology available. 4. Operability and control: experimental evaluation and further simulation, the implementation of a control scheme.

The information necessary typically includes: 1. VLE, thermodynamic model 2. kinetics and heterogeneous catalyst 3. reaction equilibrium 4. solubility, miscibility, physical properties 5. packing holdup, residence time and hydrodynamics 6. mass transfer upon packing

2.5-2-1 Catalytic Distillation Thermodynamics

Catalytic distillation thermodynamics includes fields of research such as reactive azeotropes, multiple steady states and reactive residue curves. The techniques developed serve as powerful design and evaluation tools. Their usefulness lies in the fact that they deal with the fundamental aspects of a proposed or existing system which truly influence its success.

Chapter 2: Literature Review 53 Reactive azeotropes have generated considerable interest and have prompted specialist methods of residue curve construction and interpretation. According to Frey and Stichlmair (1999) the conditions required for the formation of reactive azeotropes is that the change in concentration due to distillation is totally compensated for by the change in concentration due to reaction. They deal specifically with systems approaching equilibrium where in the presence of the chosen catalyst the reaction can be considered instantaneous. In making this assumption (or posting this limitation) they were able to develop a graphical construction method based on ordinary distillation curves, stoichiometric lines of reaction and the equilibrium line. Given that the two rates of reaction and separation must be compensated they postulate that the necessary condition for the existence of a reactive azeotrope is that the stoichiometric line must be collinear with the tangent of the residue lines. Further reactive azeotropes can only exist on the equilibrium curves for systems approaching equilibrium and hence they can be pinpointed or discounted, Figure 2.03. They suggest that the method is not suitable for systems operated in the kinetic regime and suggest the construction of reactive residue curves for such systems. Reactive azeotropes are seen to be pressure sensitive, given that the equilibrium constant is a function of temperature and that column pressure directly determines the temperature distribution of the column.

Figure 2.03 Construction for graphical determination of the the loci of possible reactive azeotropes for a reaction system of the following form a+b = c , where π is

the pole point defined by xiπ=νi/Σνj. Frey and Stichlmair (1999).

Chapter 2: Literature Review 54 A useful tool developed to map the occurrence of reactive azeotropes is that of the use of transformed variables, as developed by Ung and Doherty (1995). These allow a residue curve to be constructed which reduces the number of variables to be considered through the stoichiometry of the reaction. It also allows the lever rule of standard distillation residue curve theory to be applied. In a reacting mixture of c reacting components and I inerts for a total of C components, undergoing R equilibrium reactions, the transformed compositions are defined:

⎛⎞xNx−ν T −1 X ≡ ⎜⎟ii ref (2.46) i ⎜⎟T ⎝⎠1−ν totalx ref

⎛⎞yNy−ν T −1 Y ≡ ⎜⎟ii ref (2.47) i ⎜⎟T ⎝⎠1−ν totaly ref where ν are the stoichiometric coefficients of the reaction and N is the square matrix of stoichiometric coefficients for the R reference components of the R reactions. For a single reaction the reference component can be a reactant or a product, for systems involving multiple reaction and inerts more rigorous selection criteria apply. They proved that the necessary conditions for the occurrence of reactive azeotropes are:

X ii= Yi =−− 1,..., CR 1 (2.48)

The method developed involves evaluating the equilibria in terms of the transformed variables and then applying standard methods of azeotrope detection, such as X and Y diagrams. As in the method of Frey and Stichlmair this method is suitable for reaction systems approaching equilibrium. The strength of this method lies in dealing with systems with a larger number of components, arising from multiple chemical reactions or the presence of inert components. Okasinski and Doherty (1997). apply the method of transformed variables to various reactive distillation systems of interest including that of MTBE. They suggest that it is the presence of a reactive azeotrope near the vertex of MTBE that allows for high purity MTBE to be obtained in catalytic distillation systems.

Chapter 2: Literature Review 55 Song, Huss, Doherty and Malone (1997) were able to demonstrate the existence of reactive azeotropes experimentally using the system of the esterification of acetic acid with isopropanol giving isopropyl acetate and water over Amberlyst 15. Using 30 g of catalyst in an initial volume of 306 mL and a boilup rate of 38 mL/h, they were able to approach the limiting case of phase and reaction equilibrium. They obtained a very close fit between the predicted and observed reactive azeotrope.

Beyond reactive azeotropes the occurrence of multiple steady states (MSS) has also been explored through reactive residue curve construction and reactive distillation thermodynamics. Using the transformed variables developed for evaluation of reactive azeotropes rigorous procedures for the prediction of the likelihood of multiple steady states have been developed by Güttinger and Morari (1999) & (1999). Their analysis is also equilibrium based. They define hybrid and non-hybrid reactive distillation columns. Hybrid containing mixed reactive and non-reactive zones and non-hybrid column being reactive throughout as in the case of homogeneous catalysis. They developed a method called ∞/∞ analysis for each case. The infinities refer to infinite internal flow rates and infinite number of stages. Examining the MTBE system they develop a MSS feed region in the standard molar fraction composition space. They show that the MSS phenomenon tends to disappear as the methanol content of the feed is decreased. This lead to the compromise that decreasing MeOH in the feed helps reduce the chance of MSS however leads to isobutylene dimerisation.

Venimadhaven, Buzad, Doherty and Malone (1994) developed reactive residue curves which incorporate kinetics as well as equilibrium considerations. They followed the standard construct of a residue still and incorporated homogeneously catalysed kinetics. The Damköhler number, Da, used in three phase reactor analysis was introduced as a means of developing reactive distillation residue curves, which are not equilibrium based and which do relate to column design and was defined as: Hk Da = 0min (2.49) V0 -1 where H0 (mol) is the still liquid holdup, kmin (s ) is the pseudo first order rate constant at the minimum temperature on the boiling surface (for reaction expression defined in x, mole fractions), and V0 (mol/s) is the vapour rate. The generalised differential equation

Chapter 2: Literature Review 56 for the simple residue still incorporating reaction was developed for the case where vapour rate varies in the same way holdup does, H/H0 = V/V0, and was given as:

cr cp dxi k ⎛⎞ννmn1 =−+()xyii Dax()νν iiT −⎜⎟∏∏ a m − a n (2.50) dkξ min ⎝⎠mn==11K were the ξ is warped time (defined ξ=ln(H0/H)) , νi are stoichiometric reaction coefficients, cr and cp are the number of reactant components and product components and ai are activities. For the MTBE reaction:

H + IB + MeOH ←⎯⎯⎯⎯→ MTBE (2.51) the above equation reduces to:

dxi Da =−+()xyii r()ν ii + x (2.52) dkξ min The set of equations is solved for different values of Da (Da=0, 0.12, 0.5 50..), with the vap knowledge of K = f(T), k = f(T), Pi = f(T), stepping along the lines of integration. For calculation a column pressure must be defined and different liquid compositions near the residue curve edges are chosen. The residue curves generated exhibit various stable nodes including azeotropes, equilibrium stable pinches and reactive azeotropes. In a reactive azeotrope the rate of reaction is matched by the rate of vaporisation. The multiple steady states observed in this system also arise. The stable nodes appear as Da is increased but also move and disappear. Thiel, Sundmacher and Hoffmann (1997) extend the reactive residue theory employing Da to heterogeneous catalysis kinetics. As mentioned previously they evaluated the systems of MTBE, TAME and ETBE in considerable detail.

2.5-2-2 Limiting Regimes, Damköhler and Hatta Numbers

The Damköhler number is ratio of characteristic liquid residence time (H/V) to the characteristic reaction time (1/k1), and its implications for catalytic distillation are: • Da < 0.5: Phase equilibrium control. The reaction rate is slow relative to the residence time available and the system is dominated by phase equilibrium. If phase equilibrium controls then the reactive stages will be strongly influenced by azeotropes that are present in the non-reacting mixture. For low values of Da a large

Chapter 2: Literature Review 57 residence time is necessary. Rate-based models must be used for column design in this regime as phase equilibria are unable to account for the appearance of product. • Da > 10 Chemical equilibrium control. In this case the reaction rate is fast and the reactive stages can be assumed to be at chemical equilibrium. Each reactive stage can be modelled as an equilibrium chemical reactor. • 0.5 < Da < 10 Rate limited. Rate based models are needed as both kinetic and phase equilibrium effects are important.

There is some scope to manipulate Da in the design stage by choice of hardware and its design and by manipulation of column pressure hence temperature. When combined with knowledge of reaction equilibrium, K, then further implications about the suitability of reactive distillation may be drawn: • Da Low, K Low: Slow forward reaction and fast reverse reaction. Choice of RD poor, large hold-up required. • Da Low, K High: RD can give benefit if the required tray hold-up is not too high. • Da High, K Low: RD can be useful if the product can be removed from the reaction zone quickly enough. • Da High, K High: RD is really not required but may be used to prevent side reactions say.

Generally rate based models need to be used to model the RD and here the Hatta number can be used to determine whether or not a limiting rate model can be used. This parameter was developed in reactive absorption analysis and is based on the two-film interphase mass transfer model. The Hatta number:

Dk Ha = A 1 (2.53) kaL for which the implications are : • Ha <<1: Slow Regimes.

• k1 << kLa: Slow Kinetic Regime. Rate is most sensitive towards the intrinsic rate constant, hence improvement stems from temperature and catalyst manipulation. In modelling can assume phase equilibrium and reaction kinetics only.

Chapter 2: Literature Review 58 • k1 >> kLa: Slow Diffusion Regime. The process will benefit from improvements in vapour-liquid contacting. A rate-based model can assume equilibrium reaction rate and rate of mass transfer between phases.

• k1 ≈ kLa : Mixed regime neither dominates, the model is completely rate-based and improvements can be achieved through both catalyst and contacting device improvements. • Ha ≥1: Fast Regime. The overall rate depends on the enhancement factor. Typical of fast reactive absorption where the film model is stretched by fast reaction.

Most RD applications fall in the slow kinetic regime, k1 << kLa. If both components are expected to have high concentrations on the reactive stages than the process is in the slow kinetic regime.

2.5-3 Modelling and Simulation

Modelling and Simulation have been greatly advanced for CD processes. The most recent modelling studies are carried out using commercial simulation packages such as AspenPlus, Pro/II, HYSIS, and SpeedUp, Taylor et al (2000). While in the early stages of design typically kinetics, thermodynamics, residue curves conventional and reactive and aspects of hardware design are considered, it is useful to buildup sufficient information, formulated in such manner that it can be employed in modelling and simulation before or after more extensive experimental evaluation. Thus knowledge of what can be achieved through modelling and what is required for successful process representation is valuable.

2.5-3-1 Equilibrium Based Models

The equilibrium based model for catalytic distillation assumes that the compositions in the liquid and vapour stream leaving each stage are in equilibrium with each other. A schematic representation was given in Figure 2.04. An equilibrium stage is described using the so-called MESH equations: Material balance, Equilibrium, Summation and Heat balance equations. Such equations can be written for each stage in the stripping and rectifying section as well as for the reboiler and condenser. For the reactive section the change in the number of moles of the component i due to reaction must be

Chapter 2: Literature Review 59 considered. Chemical equilibrium or kinetics can be used to calculate the extent of reaction upon each stage.

Figure 2.04 Equilibrium stage representation in most versatile form.

Generally EQ models struggle to represent real trays where mass and energy transfer occurs across the vapour liquid interface. However only the standard physical parameters and reaction kinetics are required to initiate significant probing of a prosed design through modelling.

The MTBE MSS studies mentioned of Jacobs and Krishna (1993); Nijhuis, Kerkhof and Mak (1993); Hauan, Hertzberg and Lien (1997), where conducted using the equilibrium based RADFRAC column unit available as part of AspenPlus, along with user defined heterogeneous kinetics. Venkataraman, Chan and Boston described the inside out algorithm used in RADFRAC and demonstrated the use of the unit with numerous reactive separation applications, Venkataraman et al (1990). Gonzalez, Subawalla and Fair (1997) used RADFRAC to simulate the tert-amyl alcohol process they designed with good accuracy.

When the equilibrium approach is used to model a distillation column a correction factor is often used to account for departure from equilibrium. The correction factor is referred to as efficiency. For reactive distillation it has been shown that it is difficult to extend EQ type models through the use of Murphree stage efficiency for staged columns or HETP for packed columns. The difficultly arises due to the variation in efficiency across a single stage and the variation of HETP throughout the reactive zone.

Chapter 2: Literature Review 60 The variation across a tray for homogeneous reactive distillation can be captured by models referred to as pool models, which treat each tray as a series of CSTRs or pools, Taylor et al (2000). The complexity of defining this variability can negate the use of such parameters and suggests the use of NEQ models.

2.5-3-2 Non-Equilibrium Based Models

In practice columns rarely operate under thermodynamic equilibrium conditions. Vapour-liquid equilibrium prevails only at the interface separating vapour and liquid. The separation and reaction achieved depends upon the interphase mass and heat transfer processes. Non-Equilibrium (NEQ) based Models describe these simultaneous mass and heat phenomena and account for the multi-component interactions between simultaneously diffusing species. The common conceptualisation of a NEQ stage representing catalytic distillation is presented in Figure 2.05. Resistances to interfacial mass and energy transfer are assumed to be located in thin films adjacent to the vapour- liquid interface and to the catalyst-liquid interface. Phase equilibrium relations are employed at the interface and specific models are developed to describe heterogeneous catalysis and intra-particle mass transfer.

Figure 2.05 NEQ stage representation showing mass and energy fluxes and the use of film theory.

Chapter 2: Literature Review 61 In the NEQ model, hardware design information must be specified allowing for mass transfer coefficients, interfacial areas, liquid holdups to be calculated. The NEQ model requires thermodynamic properties not only for calculation of phase equilibrium but also for calculation of driving force for mass transfer and for taking into account the effect of non-ideal component behaviour in the calculation of reaction rates and chemical equilibrium constants. In addition physical properties such as surface tension, diffusion coefficients and viscosities for calculation of mass and heat transfer coefficients and interfacial areas, Taylor and Krishna (2000). Maxwell-Stefan equations or more rigorously by Fick’s law can be employed for modelling mass transfer in mutli- component systems. The use of Maxwell-Stefan equations in conjunction with film theory has been widely popularised, Taylor and Krishna (2000).

Several approaches have been used to model mass transfer associated with heterogeneous catalysis. They include, Baur et al (2003): 1. Pseudo-homogeneous models: where the intra-particle diffusion and reaction is simplified by the use of catalyst effectiveness factors and pseudo-homogeneous rate expressions. 2. Heterogeneous models: where detailed account is taken of intra-particle diffusion and reaction within the catalyst particle. An example of which is the dusty fluid model, which treats the porous catalyst as a dust species. The model employs effective Knudsen diffusion coefficients, Higler (2000).

2.5-4 Catalytic Distillation Hardware Design

Catalytic distillation is a form of unit intensification, where the function of a reactor is combined with that of a distillation column. It is largely hardware design, which accommodates this combination and ensures its success. While a number of hardware designs have been promoted, their characterisation in terms of fluid dynamics and ability to be customised to a particular application are lacking. Taylor and Krishna (2000) pointed out their frustration that while sophisticated NEQ design modes are available, detailed information on the hydrodynamics and mass transfer parameters for the various hardware configurations they reviewed are woefully lacking in open literature. That paradoxically such information has a vital consequence for the conversion and selectivity of RD columns, Taylor et al (2000).

Chapter 2: Literature Review 62 The importance of good and variable hardware design is further stressed by the incongruous requirements of reaction and separation. On any reactive stage the requirements of the chemical reaction of high liquid holdup for maximised conversion, are not in consonance with the requirements for good separation of high interfacial area, Krishna et al (2002). The reaction calls for small particles and high loading. The separation on the other hand prefers large interfacial area between gas and liquid and low pressure drop gained by open areas. Thus the hardware design needs to balance the extent of these opposing requirements for a particular system. Given the opposing requirements the catalyst loading in such packing as Bale or KATAPAK-S is only 20 to 25 % of the column volume.

The practical design considerations behind good hardware design and or selection according to Taylor and Krishna (2000) include: 1. Installation, containment and removal of catalyst. As the catalyst deactivates it will need to be removed, regenerated and or replaced. 2. Effective contacting of liquid with catalyst particles. The hardware design must ensure good liquid distribution as maldistribution can result in hot spots and uneven catalyst aging as well as the occurrence of side reactions. Liquid distributors, wall wipers and design which imposes mixing such as criss crossing patterns and offset packing elements as well as tray crossflow. 3. Effective vapour liquid contacting in the reaction zone. If the rate of reaction is fast and the reaction is equilibrium limited then the required size of the reactive zone is strongly influenced by the effectiveness of the vapour-liquid contacting. 4. Low pressure drop through the catalytically packed reactive section. The small catalyst particles need to be packaged such that the pressure drop is practical, creating three levels of porosity type reactor. 5. Sufficient liquid holdup in the reaction zone. The liquid holdup often tends towards the higher end. An understanding of liquid holdup, residence time and residence time distribution is essential in determining conversion and selectivity. 6. Design for catalyst deactivation. The major solution to deactivation is to add more catalyst, which in tern can lead to maldistribution. Reflux and temperature can also be used to change the reaction severity as required.

Chapter 2: Literature Review 63 The major hardware design solutions, which are being currently strived for, are on line catalyst replacement and a system, which is easily characterised and adapted. As catalytic distillation is complex it would be beneficial if the hardware reduced complexity in modelling and operation as opposed to enhancing it. The current choice of hardware for catalytic distillation can be broadly classified as: 1. Catalyst Containment, here standard catalyst used for the desired reaction is packed in some form of basket or bag device. Catalyst containment can result in random or structured packing or tray modification. 2. Catalyst Alteration, the coating of catalyst onto surfaces, the manufacture of distillation type packings from catalytic material, the change of catalyst shape and processing. Catalyst alteration has largely resulted in random type packing.

2.5-4-1 Catalyst Containment

The beauty of containment lies in its simplicity and ease of comparison to kinetics evaluated under simpler, mass transfer limitation free conditions. Some applications of containment include, Taylor et al (2000): 1. Bale Packing 2. Structured Catalytic Packing, Sulzer ChemTech KATAPAK packing. 3. Cross Flow Tray modifications 4. Downcomer modifications

Bale packing and structured catalytic packing were expanded upon given the application of interest and their employment for applications considered similar:

1. Bale Packing In Bale packing the catalyst is held in segmented fibreglass bags. The fibreglass is folded over and pockets are sewn into it. The pockets are then filled with catalyst and sewn shut. These catalyst quilts are then rolled with a layer of stainless steel mesh. The mesh promotes voidage and imparts mechanical strength. The bales are arranged in a distillation in a regular circular pattern. The layers are kept out of alignment giving good support and mixing.

Chapter 2: Literature Review 64 Bale packing was patented in 1980 by Chemical Research and Licensing as part of their MTBE process patent, Smith (1980). Their capacity and liquid holdup have been studied, Akbarnejad et al (2000), Subawalla et al (1997). Akbarnejad et al (2000) used an arrangement of seven 40 mm diameter bales per layer with a total of four layers and a total height of 800 mm in a column of 100 mm diameter. They did not employ wall wipers but offset layers from each other. They observed a distinct loading point in liquid and gas flowrates and found holdup independent of gas flow rate below the loading point. The mass transfer efficiency has also been determined. For the systems acetone/MEK and cyclohexane/n-Heptane system at 138 kPa in a 53 mm column, under total reflux, Subawalla et al (1997), determined HETP, Table 2.13. They achieved a catalyst loading of 107 kg/m3 of column with bales of 46 mm diameter and 313 mm height. The segmented bag was 166 mm long and 290 mm height Their experiments were conducted using six such bales. Each bale was equipped with three wall wipers.

Table 2.13 Approximate values of HETP over Bale packing (1997).

3 0.5 F factor (m/s.(kg/m ) HETP (m) acetone/MEK HETP (m), cC6/C7 1.25 0.55 - 2.1 0.3 - 0.9 - 0.5 1.8 - 0.37

Bale packing has been used for the manufacture of MTBE and in evaluation of a TAA process. Baur and Krishna (2002) carried out a simulation-based comparison and evaluation of using Bale packing or catalytic Raschig rings for the production of TAME. The Bale packing had a void fraction of 0.76 and an area of 169 m3/m2. The Bale column was long and narrow (18.3 m, 3.3 m) while catalytic raschig rings (25 mm) produced a short fat column (10.3 m, 4.5 m). The raschig rings had better mass transfer characteristics but poorer pressure drop and hence the dimensions arrived at. The Bale packing benefited from a shortening of the reaction zone and a lengthening of the inert stripping section, showing how it’s poorer separation characteristics can easily lead to product break down in a equilibrium limited system if the reaction section is over estimated.

Chapter 2: Literature Review 65 2. Structured Catalytic Packing, Sulzer ChemTech KATAPAK packing and Koch-Glitsch KATAMAX. They consist of two pieces of rectangular crimped wire mesh sealed around the edge there by forming a pocket of the order of 1-5 cm wide between the two screens. This pocket is then filled with catalyst and sealed. These catalyst containing sandwiches are then bound together. The crimping is arranged in a cris-cross manner at 90o to each other and forms a 45o angle to the vertical. The individual sandwiches are also arranged to be crisscross to each other. The liquid thus follows an essentially crisscross flow pattern giving a good radial distribution and promoting catalyst bed penetration. Subsequent layers are arranged out of alignment and the interface gives a boost to liquid phase mixing. The catalyst occupies between 20 and 40 vol % of the column space.

Higler and Krishna (1999) and Moritz and Hasse (1999) reported experimental studies on the non-ideal flow behaviour of the liquid phase in KATAPAK-S. Noeres and Gorak (2002) reported preliminary findings of a broader investigation of the liquid phase mixing of MULTIPAK a less intricate version of this type of structured packing. In general, these forms of structured packing are difficult to construct at a laboratory scale of 45 mm. The studies referred to above used elements of 100, 70 and 100 mm respectively. Sulzer’s official packing is available in a minimum of 100 mm diameter.

2.5-4-2 Catalyst Alteration

This is a challenging field, which has seen much progress but which has been hampered by economic feasibility and the inertia of industrial practice. One of the most attractive goals of catalyst alteration is that of producing catalytic packing that acts similarly in terms of separation properties to its inert equivalent packing. If for instance Raschig rings can be made of catalytic material they can be easily combined with these in the non-reactive part of the column. Even if their surface properties are different, a degree of uniformity in the RD column is created. This in turn produces ease of process development, as a tube column can be packed quickly exactly as required. Further advantages include a high ratio of the number of active sites to reactor volume, low pressure drop and large external surface area. A major difficulty with catalyst alterations is keeping the catalyst material fresh and active. Another difficulty experienced by those

Chapter 2: Literature Review 66 working with resin catalysts is swelling and mechanical stress, which can eventually break down the catalytic packing.

Kunz and Hoffmann (1995) and Sundmacher et al (1996), developed several methods of producing Raschig rings with catalytic activity including block polymerisation and precipitation polymerisation on a glass or ceramic support. They gave a very thorough review of the methods available. In their most successful procedure they essentially polymerised an ion exchange resin equivalent to Amberlyst 15 onto a ceramic support shaped as a Raschig ring. The precipitation polymerisation was carried out in the presence of a pore forming agent and repeated several times gaining different loadings and pore volumes. All details of the process were given and activities of 0.5 to 1.5 meq H+/g were achieved.

A low-tech practical alternative is that of sintering Amberlyst particles with plastics such as polyethylene or polypropylene to produce pellets or cylinders, Matouq et al (1994). Researches under Shigeo Goto utilised this method to study many catalytic distillation systems Abella et al (1999). They achieved pellets of 8 mm diameter and activities of 3.98 eq/kg pellet compared to 4.9 eq/kg for Amberlyst-15. The loss being attributed to pore access blockage due to the binder.

2.5-5 Dynamics and Control

Dynamics, control and controllability are being approached predominantly through modelling. Sneesby and Tade (1999), through modelling of catalytic distillation for the synthesis of ETBE have been developing control strategies and investigating aspects of dynamics such as multiplicity. They have developed aspects of inferential control by mapping the relationship between temperature and purity and conversion, Sneesby et al. (1999). They were able to find the temperature locations within the column which were most responsive to process change and used these for the purpose of control.

With the advent of tools such as DynaPlus which can be utilised with the standard units of AspenPlus or SPEEDUP dynamics and stability are being considered increasingly early in the development of catalytic distillation processes. Jimenez et al (2002) investigated dynamics for the production of butyl acetate and methanol through

Chapter 2: Literature Review 67 transesterification. They were able to develop a control structure, tune the parameters of the various PID control loops chosen and evaluate the robustness of the proposed scheme. Their process was highly novel incorporating extractive distillation with catalytic distillation and yet they were already able to consider aspects of controllability.

2.6 Catalytic Extractive Distillation

The process of catalytic extractive distillation is emerging as a solution and consequence to the presence of unbreakable conventional azeotropes occurring in systems otherwise suitable for catalytic distillation. The obvious concern is that the inherent benefits of hardware/unit intensification may be diminished by the increased complexity and additional solvent recovery column potentially required.

Some conventional azeotropes cannot be removed by reaction and it may be possible to combine standard azeotrope breaking technology with reactive distillation to achieve this. Extractive distillation and azeotropic distillation are an obvious choice for incorporation into catalytic distillation for the purpose of breaking the azeotropes, which occur, in many potential systems. These include: esterification such as that for the production of ethyl acetate, methyl acetate and butyl acetate. Beyond breaking azeotropes the solvents used can provide better conditions for reaction by increasing solubility of sparsely soluble reagents or by maintaining a homogeneous liquid phase, they may also promote reaction rates by beneficially altering activity coefficients.

One classic success story in both reactive distillation alone and the combination of reactive distillation and extractive distillation is the Eastman Chemical Co methyl acetate process, Towler et al (2000). The methyl acetate is produced by esterification in the presence of H2SO4 accordingly:

H + CH3COOH + CH3OH ←⎯⎯⎯⎯→ CH3COOCH3 + H2O (2.54)

The extractive distillation component is however subtle as the entrainer used is one of the actual reactants acetic acid. It satisfies many of the requirements of a good heavy entrainer having low volatility and altering the relative volatility of methyl acetate and methanol so as to rupture the azeotrope present. A second azeotrope exists between

Chapter 2: Literature Review 68 water and methyl acetate, however this is primarily broken by reaction. It can be noted how the azeotrope between products can be reacted away but those between products and reactants persist. The process has enjoyed genuine success for 19 years. It is a true example of unit intensification as one hybrid distillation replaced an entire flow sheet of 11 major units plus all of their heat exchangers, control systems, pumps and intermediate storage tanks.

The utilisation of the reagent as an entrainer is not possible in various new applications of transesterification being developed. Trans-esterification involves the reaction of an ester with an alcohol to produce an ester with swapped alcohol components:

+ * ** H ** * RCOOR + R OH ←⎯⎯⎯⎯→ RCOOR + R OH (2.55)

Trans-esterification is likely to be a new application of catalytic distillation that will receive much attention in the near future. This is especially the case given that some esters have a much better market value than others. Given that there are two alcohols and two esters involved in this reaction system there is considerable scope for the occurrence of unfavourable azeotropes. The interest has already begun in the application of CED technology, Jimenez et al (2002a) have investigated the kinetics of trans- esterification of methyl acetate with butanol in the presence of an extractive distillation type entrainer, o-xylene, Jimenez et al (2002a). They argued that the possible side reactions of de-hydrolysis of methyl acetate and butyl acetate does not occur to any significant extent.

They have also explored the synthesis and conceptual design of a CED using a variety of AspenTech simulation packages, Jimenez et al (2002b). Their proposed process involved four columns and was determined to be uneconomical by the authors, although very good purity and yield of product was obtained with simulation. In order to simulate the process they used UNIFAC as a thermodynamic model, and literature reported properties of industrial scale Sulzer KATAPAK-S containing Amberlyst 15. They obtained an entrainer to methyl acetate ratio of 1.98. They did not consider the possible side reactions in their simulation. Some other key design elements of their CED column are given in Table 2.14.

Chapter 2: Literature Review 69

Table 2.14 CED for methyl acetate synthesis column properties. Parameter Simulation Optimum Reflux Ratio 3 Diameter 0.55 Pressure (kPa) 101 reactive stages 6 stripping stages 5 rectifying stages 30 Entrainer to MeAc (mol ratio) 1.98 BuOH/MeAc (mol ratio) 0.988

In their simulation work Jimenez et al (2002b) found no difference between the rate based and equilibrium-stage models at steady state. They found that the system was highly coupled, which led to complex dynamics when they investigated a control strategy using the software package DynaPLUS. To develop a control structure they looked at best sensor location in terms of temperature profile. As they were feeding relatively volatile components at the top, the top column temperature was insensitive to product composition. They used a reaction zone average temperature to control the reboiler duty in order to avoid catalyst deactivation. After auto-tunning their controllers they tested the robustness of the scheme. They found that “typical process upsets” did not have a strong effect on the conversion because reaction rate increased in some stages while decreasing in other stages.

Exploring improved means of producing TBA by CED may also develop tools, which can aid the development of alternative fuels. One such alternative is that of fatty acid methyl esters (FAME) which can be used as a biologically derived diesel engine fuel Gryglewicz et al (1999). Currently there is a lot of work to make fatty acid esterification over heterogeneous catalysts a continuous process. Some very early work has been done in evaluating the feasibility of using catalytic distillation. Steinigeweg and Gmehling looked at decanoic acid esterification with methanol over A-15, Steinigeweg et al (2003). This is as large a fatty acid as can be handled by A-15 due to the catalyst’s 120oC operability limit. For larger acids and alcohols the same challenges of azeotrope breaking are being faced here as with other reactive distillation system and researchers have already started looking towards azeotrope breaking technologies. For alcohols in

Chapter 2: Literature Review 70 the C2 to C4 range Dimian and coworkers looked at adding an entrainer that would form a heterogeneous azeotrope with water thus allowing it to be removed from the reaction zone, Dimian et al (2003). The example process they used was the esterification of lauric acid with 1-propanol with the use of n-propyl-acetate as entrainer and their results looked promising with substantial increase in reaction rate. It should be noted that the entrainer chosen here, n-propyl acetate, can degrade over Amberlyst 15. Although not explicitly stated in their paper it can be assumed that this reaction is to be kept in check through the use of excess alcohol.

The removal of water from any acid catalysed heterogeneous system other than hydration should see improved reaction rates, as water tends to preferentially adsorb to acid catalyst thus producing inhibition. For alcohols in the range C5 and higher it may be possible to use the alcohol itself as a heterogeneous entrainer. One such example is that of the use of 2-ethylhexanol in dodecanoic acid esterification, Omota (2003). The azeotrope formed is the lightest in the system and upon condensation and cooling the water can be decanted and the alcohol returned to the column. The development of entrainer technology is very useful to any system, which has water present as water interacts nonidealy with a very large range of organic species forming azeotropes.

2.7 Countercurrent Fixed Bed Reactors

Catalytic reactors with two or more fluid phases can be operated in three standard arrangements: downflow cocurrent, upflow cocurrent and countercurrent. The downflow cocurrent arrangement giving the trickle bed reactor has been most commonly applied and investigated. However operation in countercurrent mode may be advantageous for reversible reactions and those inhibited by one of the reaction products. Further, the countercurrent fixed bed reactor (CFBR) is encountered in catalytic distillation as the reactive zone, albeit for a different flow regime to standard operation.

2.7-1 CFBR Applications and Technology

The countercurrent mode of operation of a conventional trickle bed packed with catalyst particles of 0.5 to 3 mm particle size at velocities of industrial relevance is difficult due to flooding. This problem can be overcome by increasing the porosity of the bed or

Chapter 2: Literature Review 71 creating separate paths for the gas and liquid flow. The reduced pressure drop, obtained by separating gas and liquid flow is achieved at the expense of mass transfer efficiency, Kundu et al (2003). There is little reason why the packings developed for catalytic distillation cannot be used in a CFBR. The CFBR offers some of the benefits of the equivalent reaction zone of a catalytic distillation column. Recent investigations have shown that the countercurrent mode of operation is advantageous in providing better gas-liquid-solid contacting and overcoming pathological reaction phenomena such as reactant or product inhibition.

In the hydrodesulphurisation and hydrogenation of heavy oils the hydrogen sulphide product strongly inhibits the reactions even with the sulphur tolerant catalysts of the mixed sulphide type, Taylor et al (2000). In cocurrent operation the bulk of the H2S is generated in a small inlet part of the bed and thus exerts its inhibiting influence on the remainder of the reactor. In a counter current mode the major part of the bed operates in a H2S lean regime.

Further, the concentration gradients are more efficient in terms of reaching yet unreacted components when operating in the countercurrent mode. In a cocurrent bed the exit concentration of the gas is at its leanest and thus there is little penetration of the liquid. Thus for instance in hydrogenation the deep removal of aromatics from the oil suffers. The benefits of countercurrent operation are such that simulations have demonstrated that the volume of catalyst can be reduced from 600 to 450 m3 for the example above, Hasselt et al (1999). The CFBR of this example was devised by Van Hasselt et al was labelled a Three Levels of Porosity reactor and consisted of a cascade of baskets holding catalyst and arranged in a brick work type pattern, with spaces between and above the baskets. The different levels of porosity are, the spaces between basket, the spaces between catalyst particles and the pores space of the particles. This concept also seems to adequately describe such catalytic distillation packing as Bale packing.

Zhang, Adesina and Wainwright (UNSW(2003)) applied a CFBR for the hydration of isobutylene to TBA over Amberlyst 15. In a similar manner to Van Hasselt et al (1999) they employed a basket to hold the catalyst. They used a series of cells to model the

Chapter 2: Literature Review 72 observed rate of reaction. In the range of liquid velocities of 0.7 to 3.75×10-05 m.s-1 they observed that reaction rate remained constant. Using mass transfer models of Goto et al (1975) for conventional trickle bed reactors, and the effective diffusivity of Velo et al (1990), they were able to account for observed reaction rates. They proposed that within their single basket the catalyst was well wetted and behaved as that of a trickle bed reactor under conditions of complete wetting. The gas to liquid mass transfer occurred predominantly at the walls of the basket. Other examples of CFBR application include the AROSAT process of aromatics hydrogenation. It employs a CFBR to maximise the hydrogenation reaction which is favoured by lower temperature and higher hydrogen partial pressure, Reilly et al (1973). Goto and Smith, (1978) investigated the air oxidation of aqueous sulfur dioxide over activated carbon using a CFBR.

2.7-1-1 Comparing CFBR to TBR

The CFBR requires openness for countercurrent flow, which reduces pressure drop but also decreases interfacial area and thus mass transfer rates. For trickle bed reactors using particles of diameter 1.5 to 2 mm the gas/liquid interfacial area reaches values between 1000 and 1500 m2/m3, whereas in a CFBR using packing of 5 mm nominal diameter the value is always lower than 500 m2/m3. The bed volume fraction occupied by catalyst is 0.65 for a TBR and 0.3 to 0.35 for a CFBR using catalytic Raschig or Berl saddles.

Transfer between fluids and particles and intraparticle diffusion are less influenced by flow direction and arrangement. Of interest are transfer between phases of mass and heat. In light of this Trambouze (1990) carried out comparative calculations for CFBR and TBR adiabatic operation with plug flow conditions for: 1. simple exothermic reaction 2. reversible reaction near isothermal (eg MTBE synthesis) and exothermic (eg aromatics hydrogenation). 3. reactions in parallel eg. combined hydrogenation and desulphurisation as described by van Hasselt. 4. case 3 with an inhibiting product

He used specific conversion as a measure of goodness. The comparison is based on the amount of catalyst not reactor volume. As an indication of the potential advantage of the

Chapter 2: Literature Review 73 CFBR he found better conversions per mass of catalyst employed for each case mentioned above. The improved utilisation of catalyst arose due to better thermal use and increased concentration potential combating inhibition effects.

Alicilar, Komurcu and Guru (2002), carried out the oxidation of cyanide over a mixed bed of the polymer Delrin (dp 5 mm) and vanadium pentoxide (dp 3 mm). Using a jacketed CFBR/TBR of 56 mm internal diameter and 64 mm bed length, they achieved higher conversion for the countercurrent mode of operation compared to the cocurrent. They attributed the improved performance to increased liquid phase holdup and mixing. According to Goto and Smith (1978), the differences between countercurrent and cocurrent operation are not large except for plug flow conditions. Under these conditions the countercurrent mode of arrangement gives much higher fractions of reactant removal from the gas phase.

2.7-2 Fluid Dynamics

The CFBR is simply the heterogenous three phase equivalent of reactive absorption as catalytic distillation is to homogeneously catalysed reactive distillation. There is a great deal of literature available on the topic of two phase fluidynamics for countercurrent applications below the loading point, such as absorption and reactive absorption, over such packing as Raschig rings and Berl saddles. This work can form a basis for the understanding of new systems such as three levels of porosity reactors and operation over catalytic distillation type packing.

2.7-2-1 Liquid Phase Mixing

The most relevant correlations for Pe and H0 over Raschig ring packing were reported in

Table 2.15. It should be noted that Peu here is interstitial or packing based. The emphasis of these studies was packing evaluation as opposed to reactor characterisation. Generally experiments were conducted using two point analysis and sloppy injection, with at least two column and packing diameters. The Galileo number, Ga, was used in conjunction with packing surface area ap and dp to allow for different packing nominal diameters. It can be seen that ReG is absent from the correlations given the lack of dependence on ReG below the loading point.

Chapter 2: Literature Review 74 Table 2.15 Dispersion Model, Pe, Correlations for countercurrent operation over Raschig Rings. Authors System (mm) Range Proposed Correlations Eq. No.

0.5− 0.33 Otake & Kunugita dp 7.6, 15.2 - (2.56) Peuu=1.9Re Ga

(1958) d 50.8 0.676− 0.44 HG0 =1.29ReL aappd

0.747− 0.693 1.968 Sater & Levenspiel ceramic 20 < ReL < 300 (2.57) PeuL=19.4Re Ga a p d p or

(1966) dp 12.7 Pe =×7.58 10−03 Re 0.703 d 101 uL, L

0.51 Bennett & Goodridge ceramic 20 < Reu,L < 200 (2.58) Peuu= 0.095Re

(1970) dp 6.4, 9.5 150

0.7− 0.32 Michell & Furzer glass 10< Reu,L < 1000 (2.59) Peuu=1.0Re Ga

(1972) dp 6.4 d 50.8, 610

0.27− 0.11 Stockar and Cevey glass 10< ReL < 500 (2.60) PeuL=1.32Re Ga (1984) d 25 p d 300

8.231 Marcias-Salinas & ceramic 340 < ReG < 4066 0.5544− 1.3 (2.61) PeBL, = 24.461ReL Ga( ap d p) or Fair (2000) dp 25 90 < ReL < 237 0.6138 d 430 PeuL, = 0.0244ReL

Chapter 2: Literature Review 75 2.7-2-2 Gas Phase Mixing

Proposed correlations for gas phase dynamics over Raschig Rings are given in Table

2.16. The form of the correlations captures the influence of ReL as well as allowing for

ReG = 0 data to be included. This data provides an important reference point and is also important in terms of pressure drop considerations.

Table 2.16 Gas Phase Mixing in the presence of Liquid Flow, for Raschig Rings (RR), in reactor based parameters.

Authors RR dp Col. Proposed Correlation (mm) id mm

Sater & 12.7 101 −0.79 −0.00299ReL PeGG=×6.23Re 10 Levenspi (2.62) el (1966)

8.231 Marcias- 25.4 430 −0.8915 −0.00075ReL (2.63) PeGG=×0.0878Re 10 ()a p d p Salinas

& 340

Fair 90

2.7-3 CFBR Analysis

There are two major approaches to the experimental treatment and the evaluation of such reactors, as a differential reactor or as an integral reactor. A differential reactor is one in which we choose to consider the rate to be constant at all points within the reactor. Since rates are concentration dependent, this assumption is usually reasonable only for small conversions or for shallow small reactors. It can also be reasonable for slow reactions where the reactor can be large, or for zero-order kinetics where the composition change can be large. An integral reactor is one in which the variation of the reaction rate is so large we choose to account for these variations in the method of analysis. Since rates are concentration dependant, such large variations in rate may be expected to occur when the composition of the reactant fluid changes significantly in

Chapter 2: Literature Review 76 passing through the reactor, Levenspiel (1972). In terms of conversion up to 10 % conversion is allowed for classification as a differential reactor. In terms of experimentation within the literature the differential treatment predominates, as it is easier to run experimentally allowing for recirculation and it is easier to model in order to evaluate parameters of interest.

2.7-3-1 Basic Reactor Treatment

Fogler (2000) develops the basic mass transfer resistance approach for trickle bed reactors in much the same manner as for slurry reactors. The analysis is based upon the following steps: mass transfer of gaseous reactant from gas to bulk liquid, mass transport from bulk liquid to outer surface of catalyst particle, intraparticle diffusion, reaction at an interior site on the pore surface and diffusion of the product towards the bulk fluid.

2.7-3-2 Integral Reactor Treatment

For first order kinetics and isothermal operation the mass conversion equations for CFBR’s can be solved analytically. Numerical solutions need to be applied to more complex kinetics. Goto and Smith (1978) demonstrated the approach for first order systems using sulfur dioxide oxidation as an example. Their analysis requires an understanding of fluid dynamics particularly in light of the dispersion model. They developed expressions for calculating reactor performance in terms of gaseous reactant removal and overall effectiveness factor.

Ramachandran and Chaudhari (1983), extended the work of Goto and Smith and demonstrated the solutions for various cases including first order reaction and m and n order power law kinetics, with constant gas phase and varying gas phase concentration. Typically beyond fluid dynamics CFBR analysis requires an understanding of intrinsic kinetics and mass transfer coefficients. In their text on three phase reactors the authors list a wide range of correlations for TBRs and CFBRs.

2.7-3-3 CFBR Mass Transfer Correlations

Some existing CFBR theory comes from pollution control operation, where the removal of gaseous components may be coupled with reaction in the liquid phase such as in reactive absorption. Such processes include sour gas scrubbing. A correlation based on

Chapter 2: Literature Review 77 an extensive amount of counter current data was proposed by Onda et al for the gas to liquid mass transfer coefficient kLa:

2/3 −0.5 0.4 kaLB⎛⎞ρρ L ⎛⎞LLu⎛⎞ u L ⎜⎟= 0.0051⎜⎟⎜⎟()adpp (2.64) agwL⎝⎠μμρ⎝⎠ awL⎝⎠ D

2 3 where ap, is the external area of particles per unit volume of reactor cm /cm and the term aw is the wetted area given by:

0.1 0.2 0.2 ⎡⎤0.75 2 2 au⎛⎞STC, ⎛⎞ρ u ⎛⎞aupl ⎛⎞ρ w =−1 exp⎢⎥ − 1.45⎜⎟⎜⎟LL ⎜⎟ ⎜⎟Ll aSagSa⎢⎥⎜⎟μ ⎜⎟ ⎜⎟ (2.65) pT⎣⎦⎝⎠⎝⎠pLT⎝⎠ ⎝⎠p

There are many correlations for the liquid to solid mass transfer coefficient in cocurrent TBR and very few for countercurrent TBR. The correlations were typically developed using the method of dissolution of sparingly soluble solids, most typically benzoic acid in water or water glycerol. Conventional trickle bed reactors and many correlations of mass transfer, effectiveness, wetting, hydrodynamics and dynamic fluid regimes have been reviewed by two giants of the field M. Herskowitz and J. M. Smith (1983). One of the more general liquid to solid mass transfer correlations, using the data of several investigations is that of Dharwadkar and Sylvester :

−0.666 −0.331 ⎛⎞μL kuslL=1.637 Re ⎜⎟ (2.66) ⎝⎠ρL D

The Chilton-Colburn analogy is usually applied to calculate heat transfer coefficients, h from ki, based on the better understanding of mass transfer.

Chapter 2: Literature Review 78 2.8 Literature Cited

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Chapter 2: Literature Review 79 Baur, R., R. Taylor and R. Krishna “Dynamic behaviour of reactive distillation columns described by a nonequilibrium stage model.” Chem. Eng. Sci. 56: 2085-2102. (2001). Beckmann, A., F. Nierlich, T. Popken, D. Reusch, C. v. Scala and A. Tuchlenski “Industrial experience in scael-up of reactive distillation with examples from C4 chemistry.” Chem. Eng. Sci. 57: 1525-1530. (2002). Bennett, A. and F. Goodridge “Hydrodynamic and Mass Transfer studies in Packed Absorption Columns.” Trans. Instn. Chem. Engrs. 48: T232-T240. (1970). Berg and Yang "Separation of teriary butyl alcohol from water by azeotropic or extractive distillation", 1992, US, 5,084,142 Bowman, F. M. and J. H. Seinfeld “Atmospheric chemistry of alternate fuels and reformulated gasoline components.” Prog. Energy Combust. Sci. 21: 387-417. (1995). Bradley, P. M., J. E. Landmeyer and F. H. Chapelle “TBA biodegradation in surface- water sediments under aerobic and anaerobic conditions.” Environ. Sci. Tech. 36: 4087-4090. (2002). Bravo, J. L. and A. Pyhalahti “Investigation in a Catalytic Distillation Pilot Plant: Vapor/Liquid Equilibrium, Kinetics and Mass-Transfer Issues.” Ind. Eng. Chem. Res. 32: 2220-2225. (1993). Chenier, P. C. "Survey of Industrial Chemistry". New York, Kluwer Academic/Plenum Publishers. Clark, J. (2001). tert-Butyl Alcohol: Chemical Properties, Production and Use, Fate and Transport, Toxicology, and Detection in Groundwater and Regulatory Standards. Oxygenates in Gasoline. A. F. Diaz and D. L. Drogos, ACS Books. Deeb, R., A. Stocking, L. A. Cohen and M. C. Kavanaugh (2001). Biodegradability of Methyl tert-Butyl Ether and tert-Butyl Alcohol. Oxygenates in Gasoline. A. F. Diaz and D. L. Drogos, ACS Books. Delion, A., B. Torck and M. Hellin “Equilibrium Constant for the Liquid Phase Hydration of isobutylene over Ion-Exchange resins.” Ind. Eng. Chem. Process Des. Dev. 25: 889-893. (1986). Delion, A., B. Torck and M. Hellin “Hydration of Isopenetenes in an Acetone Environment over Ion-Exchange resin: Thermodynamic and Kinetic Analysis.” J. Catalysis 103: 177-187. (1987).

Chapter 2: Literature Review 80 Dimian, A., F. Omota and A. Bliek “Entrainer enhanced reative distillation.” Chemical Engineeering and Processing in press. (2003). Dooley, K. M., J. A. Williams, B. C. Gates and R. L. Albright “Sulfonated Poly (styrene divinylbenzene) Catalysts.” J. Catalysis 74: 361. (1982). Drogos, D. L. and A. F. Diaz (2001). Physical properties of fuel oxygenates and additives. Oxygenates in Gasoline. A. F. Diaz and D. L. Drogos, ACS Books. Fite, C., M. Iborra, J. Tejero, J. F. Izquierdo and F. Cunill “Kinetics of the liquid phase synthesis of ethyl tert-butyl ether (ETBE).” Ind. Eng. Chem. Res. 33: 581-591. (1994). Fite, C., J. Tejero, M. Iborra, F. Cunill, J. F. Izquierdo and D. Parra “The Effect of the Reaction Medium on the Kinetics of the Liquid Phase Addition of Methanol to Isobutene.” Applied Catalysis A: General 169: 165-177. (1998). Flato, J. and U. Hoffmann “Development and Start-up of a Fixed Bed Reaction Column for Manufacturing Antiknock Enhancer MTBE.” Chem. Eng. Tech. 15: 193- 201. (1992). Fogler, H. S. "Elements of Chemical Reaction Engineering" Prentice Hall PTR, New Jersey, Third Edition, (2000) Frey, T. and J. Stichlmair “Thermodynamic Fundamentals of Reactive Distillation.” Chem Eng. Tech. 22(1): 1118. (1999). Gates, B. C. and W. Rodriguez “General and Specific Acid Catalysis in Sulfonic Acid resin.” J. Catalysis 31: 27-31. (1973). Gehlawat, J. K. and M. M. Sharma “Absorption of isobutylene in aqueous solutions of sulphuric acid.” Chem. Eng. Sci. 23: 1173-1180. (1968). Gonzales, J. C., H. Subaealla and J. R. Fair “Preparation of tert-Amyl Alcohol in a Reactive Distillation Column. 2. Experimental Demonstration and Simulation of Column Characteristics.” Ind. Eng. Chem. Res. 36: 3845-3853. (1997). Gonzalez, C. J. and F. J. R. “Preparation of tert-Amyl Alcohol in a Reactive Distillation Column. 1. Reaction Kinetics, Chemical Equilibrium and Mass-Transfer Issues.” Ind. Eng. Chem. Res. 36: 3833-3844. (1997). Goto, S. and J. M. Smith “Analysis of three phase packed bed reactors.” AIChE Journal 24(2): 294-302. (1978). Goto, S. and J. M. Smith “Trickle bed reactor performance I Holdup and mass transfer effects.” AIChE Journal 21(4): 706-713. (1975).

Chapter 2: Literature Review 81 Gotze, L., O. Bailer, P. Moritz and C. V. Scala “Recative distillation with KATAPAK.” Catalysis Today 69: 201-208. (2001). Gryglewicz, S. “Rapeseed oil methyl esters preparation using heterogeneous catalysts.” Biosource Technology 70: 249-253. (1999). Gupta, V. P. and J. M. Douglas “Diffusion and Chemical Reaction in Isobutylene Hydration within Cation Exchange Resin.” AIChE Journal 13(5): 883-889. (1967). Guttinger, T. E. and M. Morari “Predicting Multiple Steady States in Equilibrium Reactive Distillation 2 Analysis of Hybrid Systems.” Ind. Eng. Chem. Res. 38: 1649-1665. (1999). Guttinger, T. E. and M. Morari “Predicting Mutiple Steady States in Equilibrium Reactive Distillation !. Analysis of Nonhybrid Systems.” Ind. Eng. Chem. Res. 38: 1633-1648. (1999). Hanika, J., J. Kolena and Q. Smejkal “Butylacetate via reactive distillation-modelling and experiment.” Chemical Engineeering Science 54: 5205-5209. (1999). Hasselt, B. W. v., P. J. M. Lebens, H. Callis, F. Kapteijn, S. T. Sie and J. A. M. C. M. v. d. Bleek “A numerical comparison of alternative three phase reactors with a conventional trickle bed reactor. The advantages of countercurrent flow for .” Chemical Engineeering Science 54: 4791-4799. (1999). Hauan, S., T. Hertzberg and K. M. Lien “Why methyl tert-butyl ether production by reactive distillation may yield multiple solutions.” Ind. Eng. Chem. Res. 34: 987-991. (1995). Hauan, S., S. M. Schrans and K. M. Lien “Dynamic Evidence of the Multiplicity Mechanism in Methyl tert-Butyl Ether Reactive Distillation.” Ind. Eng. Chem. Res. 36: 3995-3998. (1997). Heath, H. W. and B. C. Gates “Mass Transport and Reaction in Sulfonic Acid Resin Catalyst: the Dehydration of t-Butyl Alcohol.” AIChE Journal 18(2): 321-326. (1972). Herskowitz, M. and J. M. Smith “Trickle Bed Reactors: A Review.” AIChE Journal 29(1): 1-18. (1983). Higler, A., R. Krishna and R. Taylor “Non-equilibrium modelling of reactive distillation: A dusty fluid model for heterogeneously catalysed processes.” Ind. Eng. Chem. Res. 39: 1596-1607. (2000).

Chapter 2: Literature Review 82 Higler, A. P., R. Krishna, J. Ellenberger and R. Taylor “Counter-curretn operation of a structured catalytically packed-bed reactor: Liquid phase mixing and mass transfer.” Chemical Engineering Science 54: 5145-5152. (1999). Hiwale, R. S., N. V. Bhata, Y. S. Mahajan and S. M. Mahajani “Industrial Applications of Reactive Distillation: Recent Trends.” International J. Chem. Reactor Eng. 2: 1-52. (2004). Horsley "Azeotropic Data III". (1984). Huls “Production of tert-butanol and isobutylene.” Chem. Enging(December): 60. (1983). Ihm, S. K., M. J. Chung and K. Y. Park “Activity Difference between the Internal and External Groups of Macroreticular Ion Exchange Resin Catalysts in Isobutylene Hydration.” Ind. Eng. Chem. Res.(27): 41-45. (1988). Jacobs, R. and R. Krishna “Multiple Solutions in Reactive Distillation for Methyl tert- Butyl Ether Synthesis.” Ind. Eng. Chem. Res. 32: 1706-1709. (1993). Jayadeokar, S. S. and M. M. Sharma “Ion Exchange Resin Catalysed Etherification of Ethylene and Propylene Glycols with Isobutylene.” Reactive Polymers 20: 57- 67. (1993). Jensen, K. L. and R. Datta “Ethers form ethanol 2 Transition theory analysis of the kinetics of liquid phase ethyl tert-butyl ether synthsis reaction.” Ind. Eng. Chem. Res. XXX: XXX. (1996). Jimenez, L. and J. Costa-Lopez “The production of butyl acetate and methanol via reactive and extractive distillation: II process modeling, dynamic simulation and control strategy.” Ind. Eng. Chem. Res. 41: 6735-6744. (2002a). Jimenez, L., A. Garvin and J. Costa-Lopez “The production of butyl aceate and methanol via reactive and extractive distillation: I chemical equilibrium, kinetics and mass transfer issues.” Ind. Eng. Chem. Res. 41: 6663-6669. (2002b). Kato, Y., Y. T. Honda and A. Kanzawa “Kinetic measurement on the isobutene/water/tert-butyl alcohol heat pump: dehydration of tert-butyl alcohol.” Int. J. of Energy Research 20: 681-692. (1996). Kazanskii, V. S., S. G. Entelis and N. M. Chirkov Z. Fiz. Khim. 33(6): 1409-1413. (1959). Knifton, J. F. “Capture Isobutylene with Glycol.” CHEMTECH: 43-48. (1994).

Chapter 2: Literature Review 83 Knifton, J. F., J. R. Sanderson and M. E. Stockton “Tert-Butanol Dehydration to Isobutylene via Reactive Distillation.” Catalysis Letters 73(1): 55-57. (2001). Krishna, R. “Reactive separations: More ways to skin a cat.” Chem. Eng. Sci. 57: 1491- 1504. (2002). Kroper, H., K. Schlomer and H. M. Weitz Hydrocarbon processing 48(9): 195-198. (1969). Kundu, A., S. K. Bej and K. D. P. Nigam “A novel Counter current fixed bed reactor.” The Canadian J. Chem. Eng. 81: 831-837. (2003). Kunz, U. and U. Hoffmann (1995). Preparation of catalytic polymer/ceramic ionexchange packings for reactive distillation columns. Preparation of Catalyst VI. G. Poncelet, Elsevier Science B.V. Leung, P., C. Zorrilla, F. Recasens and J. M. Smith “Hydration of Isobutene in Liquid Full and Trickle Bed Reactor.” AIChE Journal 32(11): 1839-1847. (1986). Leung, P. C. “Solubilities and Enthalpies of Adsorption Of Isobutene into tert-Butyl Alcohol Water Mixtures.” J. Chem. Eng. Data 32: 169-171. (1987a). Leung, P. C., F. Recasens and J. M. Smith “Hydration of Isobutylene in a Trickle Bed Reactor.” AIChE Journal 33(6): 996-1007. (1987b). Levenspiel, O. "Chemical Reaction Engineering". Brisbane, John Wiley & Sons. (1972). Liu, F., C. Zhang, F. Huang and C. Zhang “Studies on separation of alcohols and water by extractive distillation.” Fuel Science and Technology International 11(11): 1537-1550. (1993). Lutsyk, A. I., V. N. Mochalin and V. V. Zamashchikov “Kinetics of isobutylene hydration over a wide range of medium acidity: Effect of substrate solubility.” Kinetics and Catalysis 39(4): 463-466. (1998). Malone, M. F. and M. F. Doherty “Reactive Distillation.” Ind. Eng. Chem. Res. 39: 3953-3957. (2000). Marcias-Salinas, R. and J. R. Fair “Axial Mixing in Modern Packings, Gas and Liquid Phases: II. Two Phase Flow.” AIChE Journal 46(1): 79-91. (2000). Martel, E. H., A. Schreuders and J. P. Michaux “High purity isobutylene via extraction.” Chem. Eng. and Progress 61(3 March): 77-80. (1965).

Chapter 2: Literature Review 84 Matouq, M., T. Tagawa and S. Goto “Combined process for production of methyl tert- butyl ether from tert-butyl alcohol and methanol.” J. Chem. Eng. of Japan 27(3): 302-306. (1994). Meirelles, A., S. Weiss and H. Herfurth “Ethanol Dehydration by Extractive Distillation.” J. Chem. Tech. & Biotechnology 53: 181-188. (1992). Michell, R. W. and I. A. Furzer “Mixing in Trickle Flow through Packed Beds.” The Chem. Eng. J. 4: 53-63. (1972). Mohl, K. D., A. Kienle, E. D. Gilles, P. Rapmund, K. Sundmacher and U. Hoffmann “Steady-state multiplicies in reactive distillation columns for the production of fuel ethers MTBE and TAME: theoretical analysis and experimental verification.” Chem. Eng. Sci. 54: 1029-1043. (1999). Moritz, P. and H. Hasse “Fluid dynamics in Reactive Distillation packing Katapak-S.” Chemical Engineeering Science 54: 1367-1374. (1999). Moy, D. and M. S. Rakow "High purity isobutylene reovery", 1978, US, 4,096,194 Nicol, W. “Comparing catalytic distillation to separate reaction and distillation for the production of diacetone alcohol.” Trans. I Chem. E. 81(Part A): 1026-1032. (2003). Nijhuis, S. A., F. P. J. M. Kerkhof and A. N. S. Mak “Mutiple Steady States during Reactive Distillation of Methyl tert-Butyl Ether.” Ind. Eng. Chem. Res. 32: 2767-2774. (1993). Noeres, C., A. Hoffmann and A. Gorak “Reactive distillation: Non-ideal flow behaviour of the liquid phase in structured catalytic packing.” Chem. Eng. SCi. 57: 1545- 1549. (2002). Okasinski, M. J. and M. F. Doherty “Thermodynamic Behaviour of Reactive Azeotropes.” AIChE Journal 43(9): 2227-2238. (1997). Omota, F., A. C. Dimian and A. Bliek “Fatty acid esterification by reactive distillation. Part 1. equilibrium based design.” Chemical Engineeering Science 58: 3159- 3174. (2003). Otake, T. and E. Kunugita Chem. Eng. Tokyo 22: 144. (1958). Petrus, L., R. W. DeRoo, E. J. Stamhuis and G. E. H. Joosten “Kinetics and Equilibrium of the Hydration of Propene over Strong Acid Ion Exchange Resin as Catalyst.” Chemical Engineeering Science 39(3): 433-446. (1984).

Chapter 2: Literature Review 85 Piel, W. J. and R. X. Thomas “Oxygenates for reformulated gasoline.” Hydrocarbon Processing. (1990). Podrebarac, G. G., F. T. T. Ng and G. L. Rempel “The production of diacetone alcohol with catalytic distillation, Part I: Catalytic distillation experiments.” Chem. Eng. Sci. 53: 1067-1075. (1998). Qi, Z., K. Sundmacher, E. Stein, A. Kienle and A. Kolah “Reactive Separation of isobutylene from C4 crack fractions by catalytic distillation process.” Separation and Purification Technology 26: 147-163. (2002). Ramachandran, P. A. and R. V. Chaudhari "Three Phase Catalytic Reactors". New York, Gordon and Breach Science Publishers. (1983). Rehfinger, A. and U. Hoffmann “Kinetics of Methyl Tertiary Butyl Ether Liquid Phase Synthesis Catalysed by Ion Exchange Resin I. Intrinsic Rate Expression in Liquid Phase Activities.” Chemical Engineeering Science 45(6): 1605-1617. (1990). Reilly, J. W., M. C. Sze, U. Saranto and U. Schmidt “Aromatic reduction process is commercialised.” The Oil and Gas Journal September: 66-68. (1973). Resa, J. M., M. A. Betolaza, C. Gonzalez and A. Ruiz “Isobaric vapor-liquid equilibria of acetone-propyl ether and isopropyl ether systems. Corroboration of no Reverse volatility.” Fluid Phase Equilibria 110: 205-217. (1995). Resa, J. M. and G. A. Ruiz “Experiments of extractive distillation at laboratory scale for the rupture of the azeotropic mixture acetone and isopropyl ether.” Separation and Purification Technology 18: 103-110. (2000). Sater, V. E. and O. Levenspiel “Two-Phase Flow in Packed Beds.” I & EC Fundamentals 5(1): 86-92. (1966). Schmidt, T. C., E. Morgenroth, M. Schirmer, M. Effenberger and S. B. Haderlein (2001). Use and occurence of fuel oxygenates in Europe. Oxygenates in Gasoline. A. F. Diaz and D. L. Drogos, ACS Books. Smith, L. A. J. "Process for Preparation of Tertiary Alcohols", 1991, US, 4982022 Smith, L. A. J. "System for Separation iC4 from C4 Streams", 1980, US, 4215011 Sneesby, M. G., M. O. Tade, R. Datta and T. N. Smith “ETBE synthesis via reactive distillation 1. steady state simulation and design aspects.” Ind. Eng. Chem. Res. 36: 1855-1869. (1997).

Chapter 2: Literature Review 86 Sneesby, M. G., M. O. Tade, R. Datta and T. N. Smith “ETBE synthesis via reactive distillation 2. dynamic simulation and control aspects.” Ind. Eng. Chem. Res. 36: 1870-1881. (1997). Sneesby, M. G., M. O. Tade and T. N. Smith “Two-point control of a reactive distillation column for composition and conversion.” Journal of Process Control 9: 19-31. (1999). Song, W., S. Huss, M. F. Doherty and M. F. Malone “Discovery of a Reactive Azeotrope.” Nature 388: 561-563. (1997). Steinigeweg, S. and J. Gmehling “Esterification of a fatty acid by reactive distillation.” Ind. Eng. Chem. Res. 42: 3612-3619. (2003). Stichlmair, J. G. and J. R. Herguijuela “Separation Rgions and Processes of Zeotropic and Azeotropic Ternary Distillation.” AIChE Journal 38(10): 1523-1535. (1992). Stikkers, D. E. “Octane and the environment.” The Science of the Environment 299: 37- 56. (2002). Stockar, U. v. and P. F. Cevey “Influence of the Physical Properties of the Liquid on Axial Dispersion in Packed Columns.” Ind. Eng. Chem. Process Des. Dev. 23: 717-724. (1984). Subawalla, H. and J. R. Fair “Design Guidelines for Solid-Catalysed Reactive Distillation Systems.” Ind. Eng. Chem. Res. 38: 3696-3709. (1999). Subawalla, H., J. C. Gonzalez, A. F. Seibert and F. J. R. “Capacity and Efficiency of Reactive Distillation Bale Packing: Modeling and Experimental Validation.” Ind. Eng. Chem. Res. 36: 3821-3832. (1997). Sundmacher, K. and U. Hoffmann “Development of a new catalytic distillation process for fuel ethers via a detailed nonequilibrium model.” Chemical Engineeering Science 51(10): 2359-2368. (1996). Sundmacher, K. and U. Hoffmann “Multicomponent mass and energy transfer on different lengthscales in a packed reactive distillation column for heterogeneously catalysed fuel ether production.” Chem. Eng. Sci. 49(24A): 4443-4464. (1994). Taft, R. W. “The dependence of the rate of hydration of isobutylene on the acidity function, H, and the mechanism for oelfin hydration in aqueous acids.” J. Am. Chem. Soc. 74(5): 5372-5376. (1952).

Chapter 2: Literature Review 87 Taylor, R. and R. Krishna “Modelling Reactive Distillation.” Chemical Engineeering Science 55: 5183-5229. (2000). Thiel, C., K. Sundmacher and U. Hoffmann “Residue curve maps for heterogeneously catalysed reactive distillation of fuel ethers MTBE and TAME.” Chemical Engineeering Science 52(6): 993-1005. (1997). Thiel, C., K. Sundmacher and U. Hoffmann “Synthesis of ETBE: Residue curve maps for the heterogeneously catalysed reactive distillation process.” Chem. Eng. Journal 66: 181-191. (1997). Towler, G. P. and S. J. Frey (2000). Reactive Distillation. Reactive Separation Processes. S. Kulprathipanja. Philadelphia, Taylor and Francis: Chapter 2 Trambouze, P. “Countercurrent two-phase flow fixed bed catalytic reactor.” Chem. Eng. Sci. 45(8): 2269-2275. (1990). Tuchlenski, A., A. Beckmann, D. Reusch, R. Dussel, U. Weidlich and R. Janowsky “Reactive distillation industrial applications, process design and scale-up.” Chem. Eng. Sci. 56: 387-394. (2001). Ung, S. and M. F. Doherty “Vapor-Liquid phase equilibrium in systems with multiple chemical reactions.” Chem. Eng. Sci. 50(1): 23-48. (1995). Velo, E., L. Puigjaner and F. Recasens “Inhibition by Product in the Liquid Phase Hydration of isobutene to tert-Butyl Alcohol: Kinetics and Equilibrium Studies.” Ind. Eng. Chem. Res. 27: 2224-2231. (1988). Velo, E., L. Puigjaner and F. Recasens “Intraparticle Mass Transfer in the Liquid Phase Hydration of Isobutene: Effects of Liquid Viscosity and Excess product.” Ind. Eng. Chem. Res. 29: 1485-1492. (1990). Venimadhavan, G., G. Buzad, M. F. Doherty and M. F. Malone “Effect of kinetics on residue Curve Maps for Reactive Distillation.” AIChE Journal 40(11): 1814- 1824. (1994). Venkataraman, S., W. K. Chan and J. F. Boston “Reactive distillation using Aspen Plus.” Chem Eng. Progess 86(8): 45-54. (1990). Westerp, K. R. “Multifunctional Reactors.” Chem. Eng. Sci. 47(9): 2195. (1992). Widagdo, S. and W. D. Seider “Azeotropic Distillation.” AIChE Journal 42(1): 96-130. (1996).

Chapter 2: Literature Review 88 Yang, B.-L., S.-B. Yang and R.-Q. Yao “Synthesis of Ethyl tert-Butyl Ether from tert- Butyl Alcohol and Ethanol on strong Acid-Exchange Resins.” Reactive and Functional Polymers 44: 167-175. (2000). Zhang, C. M., A. A. Adesina and M. S. Wainwright “Isobutene hydration in a countercurrent flow fixed bed reactor.” Chem. Eng. and Processing 43: 533-539. (2003). Zhang, C. M., A. A. Adesina and M. S. Wainwright “Solubility Studies of Isobutylene in Tertiary Butyl Alcohol + Water Mixtures.” J. Chem. Eng. Data 47: 1476- 1480. (2002).

Chapter 2: Literature Review 89 3. Proposed Process and Investigation

In this chapter the motivation for the proposed process is discussed as a consequence of the literature review of Chapter 2. Key features of the application of catalytic distillation for the synthesis of TBA are detailed. The reason behind the choice of the solvent ethylene glycol and Bale packing as the catalytic hardware are presented. The chapter also details the development of the proposed process based upon evaluation of the available literature, research resources available, and the key design decisions. These components formed the basis of the investigation detailed in the body of this study.

3.1 Towards Catalytic Distillation A major potential use for TBA is considered to be that of petrol oxygenate. The requirements of this application of TBA are that any process developed should result in near pure TBA with water content approaching trace amounts. Any impurities arising from the process should at worst be flammable, simple hydrocarbons. The current Australian legislation pays special attention to the levels of sulphur, benzene and organometallics, as such these species should not be present at any concentration level. Given that direct isobutylene hydration is being considered, the process should allow for the use of diluted isobutylene of at best 50 mol % made up with inert n-butene.

The hydration of isobutylene is equilibrium limited, Delion et al (1986), and thus should benefit from the application of catalytic distillation technology. The equilibrium lines obtained at 333, 353, 373 and 393 K according to the relation of Delion et al (1986) are presented in Figure 3.01. As temperature is increased the equilibrium line moves further from the TBA vertex in correspondence to the exothermic nature of the reaction. The hydration of isobutylene is also inhibited by product TBA according to Velo et al (1988). Countercurrent operation and product separation have been shown to alleviate such inhibition, Taylor and Krishna (2000). Given that the reaction is exothermic, catalytic distillation technology can utilise this heat directly for the purpose of separation achieving some level of heat integration.

Chapter 3: Proposed Process and Investigation 90

Westerp (1992) proposed that the basic properties of a system that favour the applications of catalytic distillation are:

• A match between the reaction and distillation temperature at a moderate pressure. • Fast reactions that do not require large amounts of catalyst or reaction volume. • Difference in volatility between product and at least one key reagent. • When a solid catalyst is used liquid phase reactions are preferred as the liquid would otherwise interfere with contact between gas and catalyst. • The presence of azeotropes should be carefully considered as they can be favourable or unfavourable depending on the effect they have on reactants and products.

The indicators of favourable application of catalytic distillation technology are each discussed for the synthesis of TBA by direct hydration of isobutylene over the heterogeneous catalyst Amberlyst 15.

A match between the reaction and distillation temperature at a moderate pressure was considered practical. The hydration of isobutylene over Amberlyst 15 has been successfully carried out at temperatures as low as 303 K, Velo et al (1988). The literature available clearly promotes the use of Amberlyst 15 as a favourable choice of catalyst for the hydration of isobutylene to TBA Velo et al (1988), Leung et al (1987). Amberlyst 15 has previously demonstrated chemical, thermal and mechanical stability when applied to catalytic distillation of MTBE, Flato et al (1992) and TAA, Gonzales et al (1997). The resin’s maximum operating temperature is 393 K. At atmospheric conditions this upper limit is easily maintained given that the distillative separation of the key components water and TBA span the temperature range 335 to 373 K.

Isobutylene solubility rather than catalyst loading more pronouncedly limit the rate of hydration of isobutylene. Leung et al (1987) reported a solubility of 0.203 mol/L for 3 mol/L aqueous TBA at 60 oC and atmospheric conditions. Given the low solubility the system is also likely to exhibit low miscibility upon pressurisation. Gupta and Douglas (1967) observed the formation of two liquid phases upon compression to pressures 200

Chapter 3: Proposed Process and Investigation 91 kPa in excess of isobutylene vapour pressure, however did not report the distribution of these.

There is a favourable difference in volatility between TBA and isobutylene. However, the complete order of volatility in descending order, IB/TBA/water, leaves TBA as the middle distillate between the two reactants. The favourable order of product removed from the column in bottom stream as used with catalytic distillation processes for MTBE, ETBE, TAME and TAA, is unattainable in the TBA/water/isobutylene system regardless of operating pressure used. Unlike the case of these similar systems, volatility order cannot be reversed through manipulation of process pressure. This behaviour has been determined through the analysis of vapour pressure behaviour. Antoine coefficients were sourced for water and TBA from Sinnott (1997). The coefficients for the Wagner equation for isobutylene were sourced from Thiel, Sundmacher and Hoffmann (1997). These relations did not share a common solution for any pressure in the range of 0 to 40 atm.

The hydration of isobutylene has been previously carried out in the liquid phase within three phase reactor systems. A slurry reactor, Gupta and Douglas (1967), a packed bed reactor with external IB saturation and recycle, Velo et al (1988) and a trickle bed reactor, Leung et al (1986), and countercurrent fixed bed reactor Zhang et al (UNSW (2003)), have previously been investigated. The catalyst has a strong affinity towards water and has good wettability. Commercially, (BASF Europe), the reaction is carried out in liquid phase slurry reactors at × 30 molar excess of water, giving 88 % conversion in isobutylene, Delion et al (1986).

An azeotrope exists between water and TBA at 353 K and 64.7 mol % at a pressure of 1 atm, Gmehling (1977). This azeotrope has been experimentally proven to be relatively pressure insensitive, Horsley (1984) and Gmehling (1977). This sensitivity does not disappear with decreased or increased pressure in the range 0 to 12 atm. The central composition of the azeotrope in the x-y space ensures its formation over a relatively small number of theoretical stages in either distillative direction. Given the low solubility of isobutylene and the fact that the system cannot be operated rich in isobutylene and lean in water as this would lead to excessive formation of

Chapter 3: Proposed Process and Investigation 92 diisobutylene, it is difficult to suggest that it could be possible to react away the azeotrope. Industrial experience with application of reactive distillation for the synthesis of methyl acetate system suggests that azeotropes occurring between product and reactant cannot typically be broken by reaction, Towler (2000).

Reactive azeotropes, as opposed to conventional ones, are unlikely to occur within the TBA/water/isobutylene system. The graphical method of Frey and Stichlmair (1999) was applied to the hydration of isobutylene to TBA, Figure 3.02. The residuum lines were constructed using UNIFAC as a complete set of parameters for other thermodynamic models (NRTL, Wilson) could not be found within literature. The loci of potential reactive azeotropes can only occur close to the TBA vertex where it is possible for the stoichiometric lines originating at the pole point, π, to form tangents with the residuum lines. Referring back to Figure 3.01 it should be noted that it is not possible for this loci to cross an equilibrium line, hence it is unlikely that reactive azeotropes can form in this system.

The consideration of catalytic distillation for the synthesis of TBA in terms of these indicators of favourable application of catalytic distillation technology identified two obstacles, which were considered significant:

1. The existence of a particularly stubborn azeotrope between the reactant water and product TBA. 2. The lack of solubility of isobutylene in water and their immiscibility at elevated pressures.

These limitations in turn suggested the use of a solvent, which may improve IB solubility and allow the TBA/water azeotrope to be crossed. The use of solvent within a catalytic distillation design has been previously investigated for TAA synthesis, Gonzales et al (1997) and for transesterification, Jimenez et al (2002). Considerable process improvement was attained in each application.

Chapter 3: Proposed Process and Investigation 93 Isobutylene

0.0 1.0 0.1 0.9 0.2 0.8 0.3 0.7 0.4 0.6 0.5 T4 0.5 0.6 0.4 T3 0.7 T2 0.3 0.8 0.2 T1 0.9 0.1 1.0 0.0 TBA 0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0 Water

Figure 3.01 Equilibrium lines at T1 60oC, T2 80oC, T3 100oC and T4 120 oC.

Isobutylene

0.0 1.0 π 0.1 0.9

0.2 0.8 stoichiometric

0.3 lines 0.7

0.4 0.6

0.5 0.5

0.6 0.4 residuum 0.7 lines 0.3

0.8 0.2

0.9 0.1

1.0 0.0 TBA 0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0 Water

loci of potential conventional azeotrope reactive azeotropes & distillation boundary

Figure 3.02 Determination of the likelihood of the occurrence of reactive azeotropes within the TBA system.

Chapter 3: Proposed Process and Investigation 94 3.2 Proposed Process, Solvent Selection The use of catalytic extractive distillation (CED) is proposed as the basis of a new process for the production of TBA. Screening for a suitable solvent had formed the aim of the author’s honours thesis, Safinski (1999). Solvents with an azeotrope breaking action of either heterogeneous or homogeneous azeotropic distillation were considered.

Given the necessary temperature constraints and the similar azeotropic behaviour of TBA and water, a solvent of the heterogenous azeotropic distillation type could not be identified. The additional consideration that phase splitting is inherent to heterogeneous azeotropic distillation and can occur within the column detracted from further consideration of this mechanism of separation.

The use of a homogeneous azeotropic distillation type solvent presented itself as a superior option given the similar azeotropic behaviour and the likely formation of a single homogeneous phase of relatively lower polarity, which should promote the solubility of isobutylene. Given that water is the most polar solvent the addition of a solvent could only lower overall solution polarity and improve isobutylene solubility. Homogeneous azeotropic distillation solvents of the heavy boiling type are used in the process of extractive distillation. The combination of catalytic distillation with extractive distillation leads to catalytic extractive distillation (CED).

A large number of chemicals as available in the AspenPlus databanks and VLE handbooks Gmehling (1977), were screened using residue curves and data interpretation techniques, Safinski (1999). The organic compounds, which were found to have suitable extractive distillation properties, belonged to the following chemical groups: ethers, amides, and simple and complex glycols. Elimination of potential solvents was based upon inertness, byproduct properties, cost and availablity and ease of use. Considerations of process sustainability were also made with particular focus placed upon health and inherent safety. Finding an obviously inert entrainer proved difficult due to the high reactivity of the isobutylene/TBA/water/Amberlyst 15 system. Delion et al (1986) had found occurrence of side reactions for the use of: butylcellosolve, , cyclohexanol, THFA and acetic acid. However, they noted that the extent of these depended on the concentration of water present. Gonzalez et al (1997)

Chapter 3: Proposed Process and Investigation 95 had experienced some extent of adol condensation of acetone when studying the kinetics of isopentene hydration over Amberlyst 15.

Ethers identified, such as MTBE, were rejected as it was considered highly probable that they degrade over the strongly acidic catalyst. The glycol groups, overlapped with the ether group with a large number of poly-glycols, which contain a central oxygen atom, eg, diethylene glycol, triethylene glycol etc were also rejected.

Although amides are the least reactive of the common carboxylic acid derivatives they were rejected on the grounds that they could react with water under acid catalysed conditions, Hart (1991). Water must be kept in excess in the process of isobutylene hydration to prevent oligomerisation of isobutylene to diisobutylene, triisobutylene etc, Alcantara et al (2000). The products of the hydrolysis of amides are a carboxylic acid and ammonia. Ammonia could give rise to various other reactions within the system. Additionally, health and environmental concerns were immediately raised about their use as solvents. The amides of particular interest as solvents N-methyl-formamide and N,N-dimethyl-formamide are strong irritants and teratogenic agents.

The simple glycols can react with isobutylene, Jayadeokar and Sharma (1993). Delion et al (1986) had studied isobutylene hydration in the presence of butylcellosolve (2- butoxyethanol) a primary derivative of ethylene glycol. They found that by increasing water content to 8.7 wt % of the mixture they were able to suppress the side reaction to negligible rates. It was proposed that water, which is preferentially adsorbed by Amberlyst 15, may inhibit this reaction. The high affinity of Amberlyst 15 towards water, the use of excess water in the system and the fact that ethylene glycol need compete with water for isobutylene as opposed to reacting freely (such as in the case of degradation), suggested that a system may be possible where high selectivity towards TBA is achievable.

The simple glycols were identified for further evaluation. They were graded based on relative volatility at infinity dilution defined by:

Chapter 3: Proposed Process and Investigation 96 ∞ SAT ∞ γ LL⎛⎞P α LH = ∞ ⎜⎟SAT γ HH⎝⎠P where the subscript L and H refers to light and heavy components being separated by the solvent. The activity coefficients, γ, were estimated using the UNIFAC model. The comparison was conducted at 80 oC, the temperature of the azeotrope and is reported in Table 3.01. Table 3.01 Relative volatility of water and TBA at infinite dilution for different solvents, of the glycol sereis, determined using UNIFAC. ∞ Potential Solvent αLH 2,2,4-trimethyl-1,3-pentanediol 0.757 1,6-hexanediol 0.973 2-methyl-2,4-pentanediol 0.981 2,4-pentanediol 1.157 2,3-butanediol 1.441 2-methyl-1,3-propanediol 1.442 1,2-butanediol 1.442 1,3-butanediol 1.442 1,4-butanediol 1.442 1,2-propanediol, propylene glycol 1.933 1,3-propanediol 1.933 1,2-ethanediol, ethylene glycol 2.926

Solvents resulting in a relative volatility of 1 or less between the components to be separated have the ability to reverse the natural tendency of separation. Solvents giving relative volatilities between 1 and 1.5 have only a marginal ability to improve separation. Those giving relative volatilities above 1.5 provide substantial promotion of natural separation tendency. Of the glycol series ethylene glycol was the most promising of solvents promoting natural volatility and 2,2,4-trimethyl-1,3-pentanediol was the most promising as solvent for reversal of volatility. The natural tendency of the TBA/water system is for TBA of lower boiling point to be of higher proportion in the distillate. Thus the choice of solvent can lead to two very different designs: • CED Design A. Reversal of natural tendency of separation with TBA withdrawn as the bottoms of the CED column, Figure 3.03. • CED Design B. Enhancement of natural tendency of separation between water and TBA with TBA withdrawn at the top of the CED column, Figure 3.04.

Chapter 3: Proposed Process and Investigation 97 In Design A. the solvent acts to preferentially absorb TBA and pull it down out of a reaction zone potentially located at the top of the column. The flow sheet for Design A. where TBA is produced as bottoms was presented in Figure 3.03. Here the reaction zone, packing containing Amberlyst 15, is located in the top half of the column. The use of the entrainer here reverses the order of volatility as achieved through the use of pressure for MTBE and ETBE processes. This would be advantageous as then both reactants would be relatively more volatile than the product creating a clear path of separation. However, this advantage would be gained through considerable solvent use, given the wide potential of the natural tendency as indicated by the large difference in boiling points of 18 K and the heuristics developed for extractive distillation, Stichlmair and Fray (1992).

In Design B. the solvent acts to preferentially absorb water and pull it down into a reactive zone located towards the bottom of the column and allow pure TBA to be removed at the top of the column. The proposed flow sheet of CED for the synthesis of TBA by Design B is given in Figure 3.04. Column 1 serves as the catalytic and extractive column, while Column 2 allows for the recycle of pure solvent and unreacted water back to Column 1. The reaction zone occupies a significant portion of the Column 1 giving a rectifying section on top and a very short stripping section below. Unreacted water is returned as a saturated vapour negating the necessity of a condenser for Column 2 and utilising some portion of the duty of Column 2. The amount of water required to be recycled will depend on the required excess of water within the reaction zone and whether near pure solvent bottoms of the first column are permissible in terms of solvent boiling and solvent inertness. If the solvent is stable then the bottom of Column 1 can be relatively rich in solvent and the required reboiler duty of Column 2 can be reduced. Given that there is a strong parallelism between extractive distillation and absorption the extent of reflux should be quite small. This means that little TBA should be recycled to the reaction zone.

Chapter 3: Proposed Process and Investigation 98

UNREACTED IB / INERTS

WATER WATER

SOLVENT RECYCLE C1.

Reactive Section

TBA PRODUCT TBA & IB / INERT SOLVENT

C2. Stripping Section

Figure 3.03 CED Design A Reversal of Natural Separation Tendency of TBA and Water. Requiring a branched high molecular weight solvent such as 2,2,4- trimethyl-1,3-pentanediol of lower polarity than TBA.

Chapter 3: Proposed Process and Investigation 99

UNREACTED IB / INERTS

TBA PRODUCT

SOLVENT RECYCLE C1. Rectifying Section WATER

WATER WATER & Reactive Section SOLVENT C2.

IB / INERT Stripping Section

Figure 3.04 CED Design B Promotion of Natural Separation Tendency of TBA and water. Requiring a low molecular weight solvent such as ethylene glycol of polarity intermediate to that of water and TBA.

Chapter 3: Proposed Process and Investigation 100 As previously mentioned, a literature review of ethylene glycol and propylene glycol potential reactivity within the proposed system found that these solvents can react with isobutylene over acidic catalysts, Jayadeokar and Sharma (1993) and Alcantara et al (2000). At this point there was no information available suggesting that higher molecular weight and or more branched glycols would be slower or faster in terms of reaction with isobutylene in an aqueous system. However, in terms of pure glycols Jayadeokar and Sharma (1993) reported that isobutylene reacts three times faster with propylene glycol than ethylene glycol. Furthermore, they found that the equilibrium constant was ten times more favourable for the reaction of propylene glycol with isobutylene. It was suggested that glycols heavier than ethylene glycol would react faster with isobutylene due to a larger non-polar molecule portion and improved isobutylene solubility and attraction. The inherent attraction of the reactants is important, as isobutylene tends to react directly from the fluid phase of polar solutions according to Eley-Rideal type heterogeneous kinetics, as evident in the literature review of MTBE, ETBE, TAME and TAA kinetics, Section 2.5-4. Beyond a certain glycol size this effect may diminish due to steric hindrance and active site shielding.

For the purpose of this investigation, CED Design B, Figure 3.04, using ethylene glycol as solvent promoting natural separation tendency was chosen over Design A. for the following reasons: • Given the low solubility of isobutylene in the system, hence the low equilibrium conversion of water to TBA it was considered important to remove the TBA into the vapour phase and out of the reaction zone. Further, should selectivity become an issue then removing TBA to the vapour phase may allow for improved selectivity. • Concern that the solvent usage of Design B would be too high leading to a very large column and high solvent regeneration cost. • Given the inertness of glycols was questionable lower solvent concentrations would be preferred. • The high amount of solvent needed for Design B would result in excessive dilution and potentially further slow the reaction thus requiring increased column size through the necessity of increased catalyst loading. • The solvents giving rise to Design B such as ethylene glycol were more readily available, well characterised and significantly cheaper than those for Design A.

Chapter 3: Proposed Process and Investigation 101 The solvent ethylene glycol does not reverse the volatility order of the system isobutylene/TBA/water. The solvent should not affect the suitability of Amberlyst 15 in terms of its maximum operating temperature, as the action of an extractive distillation solvent is akin to that of an absorption solvent. The temperature profile can be controlled to some extent with the temperature of the solvent feed, while allowing for atmospheric pressure to be used. The effect of the solvent upon the reaction network and the suitability of the ion exchange capacity of Amberlyst 15 given the likelihood of side reactions occurring has yet to be determined.

3.3 Selection of Catalytic Distillation Reactor At present a hardware selection guide is not available in open literature and a generalised selection method has yet to be developed. Baur and Krishna (2002) carried out the basic design formulation for the synthesis of TAME comparing catalytic Bale and Raschig Ring packing. Each form of packing, structured or random, was capable of achieving a successful design.

Of the possible hardware choices outlined in Section 2.5-4 bale packing was chosen for this project for the following reasons: • Bale packing is an established industrially significant form of catalytic distillation hardware. • The synthesis of tert-amyl-alcohol TAA in the presence of large amounts of the solvent acetone had been previously conducted over Bale packing containing Amberlyst 15, with a considerable measure of success being reported, Gonzales et al (1997). This gave confidence to the ability of Bale packing to accommodate both the requirements of reaction and separation in the presence of a solvent. • Given the existing laboratory equipment, the packing had to be either random or structured as opposed to tray column alterations. • Bale packing does not affect inherent catalyst activity such as catalytic material coating, forming and pelletising with a binder Kunz et al (1995). • Bale packing can be easily fabricated and its design was relatively well defined in the open literature. Structured packing such as that of KATAPAK-S was relatively poorly defined and unavailable commercially at the required small pilot scale size of 45 mm diameter column.

Chapter 3: Proposed Process and Investigation 102 • Bale packing is constructed with relatively low skill procedures. • Open literature was available on the topics of capacity and holdup for small (d 45 mm) and pilot scale columns, Akbarnejad et al (2000), Xu et al (1997), and Subawalla et al (1997). The holdups measured were of the magnitude required for catalytic distillation. • Open literature was available on the topic of mass transfer, Subawalla et al (1997), and suitable mass transfer correlations had also been developed, Zheng and Xu (1992).

3.4 Proposed Investigation In summary, the following key design decisions regarding the application of catalytic extractive distillation were made prior to further investigation:

• Direct hydration of isobutylene as a pathway to TBA. • Atmospheric conditions for column pressure. • Amberlyst 15 as catalyst. • Ethylene glycol as process solvent in accordance to CED Design B. • Extractive distillation and reaction in the same section of the column as a consequence of CED Design B. • Bale packing as catalytic hardware.

Considering the evaluation of the process summarised previously, the following tasks were defined in order to accomplish the objectives of the study:

1. Determine the solubility of isobutylene within the isobutylene/water/ethylene glycol/TBA system. Judge the effectiveness of ethylene glycol in terms of isobutylene solubility improvement. 2. Determine the reactivity of ethylene glycol within the system. Determine the relative rates of isobutylene hydration and glycolisation allowing for optimisation for TBA production. Determine a kinetic model for the hydration of isobutylene in the presence of ethylene glycol over Amberlyst 15.

Chapter 3: Proposed Process and Investigation 103 3. Fabricate a relevant form of catalytic distillation hardware such as Bale packing allowing for countercurrent operation at a laboratory scale. Evaluate the fluid dynamics and hydraulic capacity of the packing obtained. Improve and develop the hardware to suit the requirements of the proposed process. 4. Evaluate the effectiveness of Bale packing in terms of the key separation of extractive distillation of TBA/water using ethylene glycol. Determine experimentally the effectiveness of the solvent in breaking the azeotrope between water and TBA. 5. Combine separation and reaction over Bale packing within a small pilot scale unit and demonstrate feasibility of the proposed catalytic extractive distillation process. Establish whether or not the addition of the solvent leads to any measurable enhancement of the process in terms of TBA purity, rate of TBA production and utilisation of isobutylene. Determine the significant parameters governing the performance of the system with respect to selectivity towards TBA. 6. Characterise the transport of the process limiting reagent isobutylene from the vapour phase to the site of reaction in the presence of the solvent ethylene glycol.

In response to this list of tasks, specific investigation strategies were researched, developed and validated. The investigation carried out was largely experimental coupled with appropriate forms of analysis. Given the different nature of the tasks, as they focus on different engineering issues and required different resources and analysis, they were divided into self-contained yet sequentially related chapters.

Task 1 and 2 in determining isobutylene solubility and the kinetics of isobutylene hydration to TBA in the presence of ethylene glycol as solvent formed the subject matter of Chapter 4. These investigations were carried out in a bench top stirred basket reactor. The effects of external and internal mass transport upon reaction kinetics were evaluated. A thermodynamically consistent heterogenous kinetic model was developed based upon observed reaction rates.

Task 3 formed the basis of Chapter 5. Bale packing was prepared and characterised allowing for improved catalytic distillation development. The two-phase fluid dynamics of Bale packing under countercurrent conditions were determined using specifically

Chapter 3: Proposed Process and Investigation 104 developed tracer techniques and a high degree of automation and data acquisition. Throughout the study, the fluid dynamics of Raschig rings (8 mm) of similar porosity were concurrently evaluated. This strategy provided a literature-based benchmark for the results obtained.

Task 4 formed the basis of Chapter 6, the extractive distillation study over Bale packing. The aim of this study was to determine the requirements of separation, so that they could be compared with and assessed against those of reaction. A simplified, semi- batch mode of operation as first proposed by Resa et al (2000) was adopted, leading to a reduction of variables and a system suitable for the catalytic distillation phase of the study.

The fulfilment of tasks 5 and 6 formed the subject of Chapter 7. Given the process- limiting nature of isobutylene, its mass transport over Bale packing was determined under countercurrent fixed bed reactor conditions in the presence of the solvent. The strategy for using catalytic distillation packing as a fixed bed reactor for the independent analysis of mass transfer has been previously adopted by Popken et al (1999). The combined action of reaction and distillative separation were studied and the effectiveness of ethylene glycol evaluated.

Observations and conclusions made throughout the completion of these studies suggested in a return to Task 3 and improvement to the catalytic distillation reactor design. In Chapter 8 a new design was proposed and investigated. Conclusions and recommendations are presented in Chapter 9.

3.5 Literature Cited Akbarnejad, M. M., A. A. Safekordi and S. Zarrinpashne “A study on the capacity of reactive distillation bale packings: experimental measurement, evaluation of existing models and preparation of a new model.” Ind. Eng. Chem. Res. 39: 3051-3058. (2000). Alcantara, R., E. Alcantara, L. Canoira and M. J. Franco “Trimerization of isobutylene over Amberlyst 15 catalyst.” Reactive and Functional Polymers 45: 19-27. (2000).

Chapter 3: Proposed Process and Investigation 105 Alcantara, R., L. Canoira, C. Fernandez-Martin, M. J. Franco and J. I. Martinez-Silva “Sythesis of tert-Butoxy-2-Propanol (PGTBE) from Propylene Glycol and Isobutylene in a Packed Trickle-Bed Reactor on Acid Catalyst.” Reactive and Functional Polymers 43: 97-104. (2000). Baur, R. and R. Krishna “Hardware selection and design aspects for reactive distillaition columns. A case study on synthesis of TAME.” Chemical Engineeering and Processing 41: 445-462. (2002). Delion, A., B. Torck and M. Hellin “Equilibrium Constant for the Liquid Phase Hydration of isobutylene over Ion-Exchange resins.” Ind. Eng. Chem. Process Des. Dev. 25: 889-893. (1986). Flato, J. and U. Hoffmann “Development and Start-up of a Fixed Bed Reaction Column for Manufacturing Antiknock Enhancer MTBE.” Chem. Eng. Tech. 15: 193- 201. (1992). Frey, T. and J. Stichlmair “Thermodynamic Fundamentals of Reactive Distillation.” Chem Eng. Tech. 22(1): 1118. (1999). Gmehling, J. and U. Onken "Vapour Liquid Equilibrium Data Collection, Aqueous Organic Systems". Frankfurt, DECHEMA. (1977). Gonzales, J. C., H. Subaealla and J. R. Fair “Preparation of tert-Amyl Alcohol in a Reactive Distillation Column. 2. Experimental Demonstration and Simulation of Column Characteristics.” Ind. Eng. Chem. Res. 36: 3845-3853. (1997). Gonzalez, C. J. and F. J. R. “Preparation of tert-Amyl Alcohol in a Reactive Distillation Column. 1. Reaction Kinetics, Chemical Equilibrium and Mass-Transfer Issues.” Ind. Eng. Chem. Res. 36: 3833-3844. (1997). Gupta, V. P. and J. M. Douglas “Diffusion and Chemical Reaction in Isobutylene Hydration within Cation Exchange Resin.” AIChE Journal 13(5): 883-889. (1967). Hart, H. "Organic Chemistry A Short Course". Melbourne, Houghton Mifflin Company. (1991). Horsley "Azeotropic Data III". (1984). Jayadeokar, S. S. and M. M. Sharma “Ion Exchange Resin Catalysed Etherification of Ethylene and Propylene Glycols with Isobutylene.” Reactive Polymers 20: 57- 67. (1993).

Chapter 3: Proposed Process and Investigation 106 Jimenez, L. and J. Costa-Lopez “The production of butyl acetate and methanol via reactive and extractive distillation: II process modeling, dynamic simulation and control strategy.” Ind. Eng. Chem. Res. 41: 6735-6744. (2002). Kunz, U. and U. Hoffmann (1995). Preparation of catalytic polymer/ceramic ionexchange packings for reactive distillation columns. Preparation of Catalyst VI. G. Poncelet, Elsevier Science B.V. Leung, P., C. Zorrilla, F. Recasens and J. M. Smith “Hydration of Isobutene in Liquid Full and Trickle Bed Reactor.” AIChE Journal 32(11): 1839-1847. (1986). Leung, P. C. “Solubilities and Enthalpies of Adsorption Of Isobutene into tert-Butyl Alcohol Water Mixtures.” J. Chem. Eng. Data 32: 169-171. (1987). Leung, P. C., F. Recasens and J. M. Smith “Hydration of Isobutylene in a Trickle Bed Reactor.” AIChE Journal 33(6): 996-1007. (1987). Popken, t., R. Geisler, L. Gotze, A. Brehm, P. Moritz and J. Gmehling “Reaction kinetics and Reactive Distillation-On the transfer of Kinetic Data from a Batch Reactor to a Trickle bed Reactor.” Chem. Eng. Tech. 21(5): 401-404. (1999). Resa, J. M. and G. A. Ruiz “Experiments of extractive distillation at laboratory scale for the rupture of the azeotropic mixture acetone and isopropyl ether.” Separation and Purification Technology 18: 103-110. (2000). Safinski, T., 2-Methyl-1,3-propanediol as a cosolvent towards the catalytic distillation of tert-butyl alcohol, Honours, UNSW, Sydney, Australia, 1999 Sinnott, R. K. "Coulson and Richardson's Chemical Engineering". Sydney, Butterworth and Heinemann. (1997). Stichlmair, J. G. and J. R. Herguijuela “Separation Rgions and Processes of Zeotropic and Azeotropic Ternary Distillation.” AIChE Journal 38(10): 1523-1535. (1992). Subawalla, H., J. C. Gonzalez, A. F. Seibert and F. J. R. “Capacity and Efficiency of Reactive Distillation Bale Packing: Modeling and Experimental Validation.” Ind. Eng. Chem. Res. 36: 3821-3832. (1997). Taylor, R. and R. Krishna “Modelling Reactive Distillation.” Chemical Engineeering Science 55: 5183-5229. (2000). Thiel, C., K. Sundmacher and U. Hoffmann “Synthesis of ETBE: Residue curve maps for the heterogeneously catalysed reactive distillation process.” Chem. Eng. Journal 66: 181-191. (1997).

Chapter 3: Proposed Process and Investigation 107 Towler, G. P. and S. J. Frey (2000). Reactive Distillation. Reactive Separation Processes. S. Kulprathipanja. Philadelphia, Taylor and Francis: Chapter 2. Velo, E., L. Puigjaner and F. Recasens “Inhibition by Product in the Liquid Phase Hydration of isobutene to tert-Butyl Alcohol: Kinetics and Equilibrium Studies.” Ind. Eng. Chem. Res. 27: 2224-2231. (1988). Westerp, K. R. “Multifunctional Reactors.” Chem. Eng. Sci. 47(9): 2195. (1992). Xu, X., Z. Zhihai and S. Tian “Study on Catalytic Distillation Processes. Part III: Prediction of Pressure Drop and Holdup in Catalyst Bed.” Trans. I Chem. E. 75: 625. (1997). Zhang, C. M., A. A. Adesina and M. S. Wainwright “Isobutene hydration in a countercurrent flow fixed bed reactor.” Chem. Eng. and Processing 43: 533-539. (2003). Zheng, Y. and X. Xu “Study on Catalytic Distillation Processes: Part I: Mass Transfer Characteristics in Catalyst Bed within Column.” Trans. I Chem. E. 70: 459. (1992).

Chapter 3: Proposed Process and Investigation 108 4. Kinetics of the Hydration of Isobutylene to TBA over Amberlyst-15 in the Presence of the Solvent Ethylene Glycol

The hydration of isobutylene to tert-butyl alcohol (TBA) over Amberlyst 15 has been examined previously in this study (cf. Section 2.3-3). However, relatively few studies have focused on the hydration in the presence of a solvent and in particular ethylene glycol. Thus, the first experimental step in the development of catalytic extractive distillation in the presence of ethylene glycol is considered to be that of gaining an understanding of how the solvent affects kinetics. Further, the literature review and considerations of fundamental design suggest that it is difficult to identify a solvent, which would remain entirely inert in this highly reactive system. Ethylene glycol can react with isobutylene to form ethylene glycol mono-tert-butyl ether (MET) therefore it is necessary to understand the conditions favouring its selectivity, such that the reaction network established may be characterised and optimised for the desired product TBA. Thus it was the ultimate purpose of this work to quantify the effects of temperature; and reagent and solvent concentration upon reaction rate. Additionally the aim has been to develop kinetic expressions free of external and internal mass transfer resistances that could be used in further design and reactor analysis work.

The benefits of the use of ethylene glycol as a solvent for catalytic distillation for the synthesis of TBA include: improved isobutylene solubility and the enhancement of the relative volatility of water and TBA allowing for distillation boundaries to be crossed. In order to analyse the liquid phase reaction kinetics and test the first potential benefit the solubility of isobutylene required characterisation. A complete understanding of the behaviour of isobutylene solubility in mixtures of water, ethylene glycol and TBA was considered important as it is the limiting reagent in terms of both reaction rate and process viability. Isobutylene needs to be the limiting reagent, as an isobutylene rich system cannot be operated due to the high likelihood of isobutylene polymerisation.

Chapter 4: Kinetics of Hydration.... 109 Thus it was an additional aim of this work to quantify and model isobutylene solubility in a mutlicomponent medium.

The experimental investigation was carried out in a semi-batch rotating basket reactor, batch in liquid and continuous in gas phase. In the initial stages of this work substantial focus was placed on reactor design, characterisation and method development. As such, an additional goal of the work was critical evaluation of the method for this specific system. It should be clarified that this work was in no way exhaustive and that it was not the purpose of this work to generate mechanistic information regarding the complete reaction network.

This chapter outlines the theoretical background regarding practical aspects of heterogenous catalysis. Details of the method and decisions made in its development are also outlined. Presentation of the results and associated discusses are provided with respect to the literature available. Firstly, in relation to the runs in the absence of solvent and then, with solvent presence. Pseudo homogenous models are considered throughout the discussion allowing for prominent effects to be evaluated. Heterogenous kinetics models, their development and evaluation are described lastly for the hydration of isobutylene exclusively.

Chapter 4: Kinetics of Hydration.... 110 4.1 Background Theory . 4.1-1 Mass Transport The general three-phase reaction system can be represented by:

A + vB ⎯⎯→ Products (4.01) where A is a reactant present in the gas phase, B is a non-volatile reactant present in the liquid phase and v is a stoichiometric parameter. The component A is involved in the following basic steps, Fogler (1999):

1. Absorption form the gas phase into the liquid phase at the bubble surface 2. Diffusion in the liquid phase from the bubble surface to the bulk liquid 3. Diffusion from the bulk liquid to the external surface of the solid catalyst 4. Internal diffusion of the reactant in the porous catalyst 5. Reaction within the porous catalyst

And the governing mass transfer equations are defined:

Step 1. CPHii= (4.02)

Step 2. RAbbib=−ka() C C (4.03)

Step 3. RAcpbs=−ka() C C (4.04)

Step 4. & 5. RAm=−η () rAs (4.05)

The film theory model makes use of an equilibrium relationship at the interface. For sparingly soluble Henry’s law defined:

PyPHxAA= = A (4.06)

The value ap, the particle/packing specific area and for packed bed reactors can be calculated by: 6(1−ε ) ap, packed = (4.07) d p

Chapter 4: Kinetics of Hydration.... 111 where ε is the bed voidage, dp is the equivalent diameter. The rate of mass transfer of A from the gas to liquid can be more generally expressed as:

* RALbb=−Ka( A A) (4.08) where A* is the concentration of A in the liquid in equilibrium with the gas phase. The

KLab term is an overall coefficient and can be broken down to the mass transfer coefficients found above: 111 =+ (4.09) KaLb ka Lb Hka gb

If component B is in excess compared with the dissolved concentration of A (Bb >> A*), then a mass transfer limitation for B does not exist. The concentration of B in the catalyst then is the same as that existing in the bulk liquid. If this condition is not satisfied then the additional steps of B must be accounted for in a similar manner as for A.

The solid catalyst is usually porous and thus the components A and B will experience intraparticle diffusion and some resistance to mass transfer. The effect of these intraparticle gradients has been characterised by the catalyst effectiveness factor defined as: actual rate of reaction r η ==A, observed (4.10) rate of reaction without diffusional limitation rAS

In the absence of mass transfer limitation the observed rate rA,observed would be equal to the rate obtained if the concentration within the catalyst particle was equal to that at its surface, CAS. For a spherical catalyst particle (with first order kinetics):

11⎡ ⎤ ηφ=−⎢coth 3 ⎥ (4.11) φ ⎣ 3φ ⎦ where φ is the Thiele modulus. For the reactions that are first order with respect to A and zero order with respect to B it is defined as:

Chapter 4: Kinetics of Hydration.... 112 0.5 R ⎡ ρk1 ⎤ φ1 = ⎢ ⎥ (4.12) 3 ⎣ De ⎦

where De is the intraparticle effective diffusivity. The effective diffusivity can be estimated according to: Dε D = cat (4.13) e τ

where εcat is the catalyst particle porosity and τ is the pore tortuosity, a value generally in the range 1 to 6 is expected, Chaudhari et al (1980). If a commercial catalyst is being used these values may have been previously determined. The diffusivity D can be estimated using the Wilke and Chang equation, given in Hines and Maddox (1985). If

φ1 < 0.2 then the value of η can be taken as unity and if φ1 > 5 then η can be taken as

1/φ1. The internal mass transfer characteristics of the hydration of isobutylene have been examined quite extensively, Velo et al (1990), Ihm et al (1988) Gupta et al (1967), and thus more informed estimates can be made.

There are two broad approaches to determining internal and external mass transport-free kinetics that of:

1. Eliminating mass transport limitation such that the reaction rate is rate determining most commonly achieved through the use of small particles, agitiation and or high velocities. 2. The characterisation of transport limitation through the evaluation of the effectiveness factor, most commonly through the use of at least two sufficiently different particle sizes, Safinski and Adesina (2005).

For a multiphase reactor, Mears (1971) has suggested that external mass transfer would be negligible if: −rRnρ Aobs bulk < 0.15 (4.14) kCc Abulk where n = reaction order R = catalyst particle radius, (m)

Chapter 4: Kinetics of Hydration.... 113 -3 ρbulk = bulk density of catalyst (kg.m ) -3 CAbulk = bulk concentration of A, (kmol.m ) -1 kc = mass transfer coefficient, (m.s )

Equation (4.15) suggests that external mass transfer in heterogeneous catalysis may be improved through:

• Increasing the mass transfer coefficient kc, which is most readily achieved through increasing the fluid velocity around the catalyst

• Increasing the bulk concentration of A, CAbulk

The Weisz-Prater modulus Φ (observable Thiele modulus) was developed to allow investigators of kinetics to determine whether internal mass transfer was limiting prior to having a complete knowledge of kinetics and the rate constant “k”. It proceeds from the Thiele modulus and can be applied as a first test for the absence of internal mass transfer limitation and is defined: −rRρ 2 Φ= A, obs catalyst (4.15) DCeAS where R = catalyst particle radius, (m) -3 ρcatalyst = particle density of catalyst (kg.m ) -3 CAS = concentration of A at the surface, (kmol.m ) -2 De = effective diffusivity, (m.s )

For first order kinetics:

2 Φ=φ1 η (4.16)

The following conclusions can be drawn from the use of the Weisz modulus: Φ << 1 : then η ≈ 1, and internal mass transfer limitations are insignificant. Φ >> 1 : then η is substantially < 1 and internal mass transfer limitations are significant.

Upon closer inspection it can be noted that improvement towards the elimination of internal mass transfer can be made by:

Chapter 4: Kinetics of Hydration.... 114 • Reducing the size of the catalyst • Reducing the observed reaction rate through reduction in catalyst loading and reactant concentration

• Increasing effective diffusivity, De

• Increasing the surface concentration CAS via increased partial pressure in A

As Φ is proportional to the square of catalyst size, reducing the catalyst size has a more dramatic effect upon Φ than changes in any other variable. As the catalyst size is reduced, intuitively more of the reaction occurs at its outer surface hence insignificant concentration profiles can be established towards the centre of the particle. In the case of the hydration of isobutylene over Amberlyst 15, De is known to decrease with temperature.

4.1-2 Kinetic Models Mth and nth order pseudo-homogeneous power law reaction models can most easily describe the reaction kinetics:

k1 aCA + bCB ←⎯⎯⎯⎯→ cCC (4.17) k−1

mn o −=rkCCkCAABC11 −− (4.18)

These simple models become appreciated when attempting to carry out any general reactor design calculations or simple evaluation of potential reactors.

The mechanistic models describing kinetics over heterogeneous catalysts most typically found in literature describe the intrinsic reaction rate in terms of surface phenomena. At the surface the following successive steps occur: 1. Adsorption onto an available active site. Inert components, other reacting components and or products may occupy the sites. 2. Reaction with a molecule on an adjacent site of the same or different type, or reaction with a molecule which has not adsorbed to a site, or decomposition of the adsorbed molecule. 3. Products are desorbed from the surface, which then frees the site.

Chapter 4: Kinetics of Hydration.... 115 All components are assumed to be in equilibrium. The rates derived from the postulated mechanisms are of the form:

()(kinetic term driving force or displacement from equilibrium) −=r (4.19) A ()resistance term

For the reaction system:

k1 aCA + bCB ←⎯⎯⎯⎯→ cCC (4.20) k−1 the basic models with the inclusion of an inert adsorbing component D, which can be a gas diluent or a solvent, in terms of concentrations Ci and for surface reaction taken as the limiting step, are as follows:

Langmuir-Hinshelwood (LHHW) kinetics:

⎛⎞CC kK1 AB K⎜⎟ CC AB− ⎝⎠K r = n (4.21) ()1++++KCAA KC BB KC CC KC DD

Eley-Rideal (ER) kinetics with B adsorbed and A not:

⎛⎞CC kK1 BAB⎜⎟ CC − ⎝⎠K r = n (4.22) ()1+++KCBB KC CC KC DD

These fundamental models can be adapted to fine-tune the correlation between the model and the observed reaction rate data and potential infer more about the underlying mechanism over the heterogenous catalyst. Successive adsorption terms can be dropped,

KDCD, KCCC etc implying that these species do not adsorb on the surface or their influence in reaction inhibition is minor. The exponent n may be assigned specific values suggesting the number of reagent involved in the rate-determining step. The “1” in the denominator can be dropped implying that the fraction of the free accessible sites is negligible under the reaction conditions.

Chapter 4: Kinetics of Hydration.... 116 These models have the implicit assumption that the distribution coefficient, λ, between bulk and catalyst phase, for reactants and products is constant, independent of temperature and concentration. As in the case of resin catalysts of gel and macroporous type the distribution of species may be skewed for components with higher affinity for the resin which cause swelling. A power law model can be employed to account for such behaviour:

m CCAcat..= λ Abulk (4.23)

The substitution of which in Equation (4.22) or (4.23) results in the incorporation of λ in the other constants and the persistence of component exponents. Gonzalez and Fair (1997) used such an empirical approach in modelling the heterogeneous kinetics of the hydration of isoamylene over Amberlyst 15 in the presence of the solvent acetone.

The parameters obtained have to be accessed for physiochemical meaning. The rate coefficients should increase with temperature and the equilibrium constant change appropriate for an endothermic or exothermic reaction. The equilibrium constant of adsorption values should be positive. The activation energy calculated should be positive and of the order of magnitude suited to the catalyst studied Parra et al (1995).

If LHHW type heterogeneous kinetics apply, with single step rate control and equilibrium chemisorption for other than the rate determining step, then use can be made of thermodynamic criteria to determine the legitimacy of the derived kinetic adsorption coefficients in terms of adsorption enthalpy, ΔH and entropy ΔS. These are obtained experimentally from the plot of, Adesina (1988):

ΔHS1 Δ ln K =+AD AD (4.24) AD R TR

In terms of enthalpy considerations the occurrence of endothermic chemisorption is suggested by Carberry (1976) as extremely rare, hence K should not increase with T.

Chapter 4: Kinetics of Hydration.... 117 Boudart, Mears and Vannice (1967) suggested that for a mechanism to be valid the entropy of chemisorption should also be considered. They suggested the following rules: 1. ΔS < 0 as the adsorbed state is more ordered than the fluid state.

2. −ΔSS < G as a molecule cannot loose more entropy than it posses

3. −ΔS =10 entropy units

4. −ΔSH <12.2 − 0.0014 Δ

where rules 3 and 4 were derived empirically based upon a large volume of literature data.

Adesina (1988) summarised and built upon improvements made to these basic rules and suggested that the experimental adsorption entropy on a Langmuirian surface should lie between two limiting cases/values: 1. the adsorbed species in a completely localised, immobilised non-interacting state,

giving ΔSloc 2. the adsorbed species is essentially free to move anywhere on the 2-dimensional

surface, giving ΔS2d

Thus, the experimental value should lie between the constrained and free as:

ΔSloc < ΔSexp < ΔS2d (4.25)

Based on statistical thermodynamics, expressions were developed to evaluate these bounds for gas-solid reactions. This improved approach was used to discriminate between an oxygenated intermediate mechanism and a carbide mechanism for CO hydrogenation to methane over group VIII metals.

Chapter 4: Kinetics of Hydration.... 118 4.2 Experimental Reliable kinetic data is of the utmost importance for the proper design and interpretation of a catalytic distillation column. Reaction kinetics need to be obtained at the conditions of the catalytic distillation reaction zone or as close to these as practicable. The kinetics were determined under atmospheric conditions as the column would be run as such. A wide range of ethylene glycol concentrations 0 to 90 mol % were used to accommodate any concentrations brought forward by the separation study. This section describes in detail equipment, procedure, experimental design and consideration made in the choice of these.

4.2-1 Stirred Basket Reactor The reactor used is illustrated in Figure 4.01. It is cylindrical, constructed of stainless steel and has a volume of 400 mL. It is fitted with a sintered metal distributor and a separate gas distribution chamber. The chamber and reactor lid and main shell are sealed with Viton o-rings. The reactor has been designed with 8 highly universal ¼” female NPT ports, allowing for reflux, sampling, temperature measurement and gas delivery. It is heated by an externally wound Isopad (328 W) heating tape.

As shown in Figure 4.01 the reactor incorporates a highly efficient sintered metal gas distributor. It had previously been operated as a bubble column slurry reactor. It was considered necessary to modify it to include stirring in order to prevent bubble coalescence beyond the distributor, given the high viscosity of the intended solvent ethylene glycol and the likelihood of external mass transfer resistances.

Two options were fabricated for the new stirred reactor a Rushton turbine and baffle system and a basket type impeller baffle system. The Rushton turbine and baffle system were designed according to the standard aspect ratios proposed and characterised by Rushton, Costich and Everett (1950), with four full length baffles used and a six blade turbine.

The impeller of the reactor was made from a stainless steel wire mesh folded into a 10 × 15 × 25 parallel pipe such that upto 10 g.L-1 of catalyst can be handled by the two

Chapter 4: Kinetics of Hydration.... 119 baskets. Zhang and Adesina (UNSW (2003)) had previously found that this is an optimum value for TBA kinetics in a slurry reactor with similar fluid dynamics. The baskets were sealed and filled by the same threaded opening used with a M5 bolt to attach them to the central holding ring. The stainless steel mesh sourced was of aperture

315 μm. Goto and Saito (1984) using Amberlyst 15 of dp = 780 μm and varying the aperture of the baskets found that below 160 μm, at various Ω (rpm), solid to liquid mass transfer was dominant. They also disproved the concern that when used with three phase reaction systems, gas bubbles may become entrained in the basket and significantly lower mass transfer rates. They found mass transfer rates to be significantly higher than in three phase stirred slurry reactors.

Following trial runs and further consideration the basket reactor was selected for the following reasons: ™ Significant catalyst attrition was observed during trial runs using the Rushton turbine. This in turn could result in: • skewed particle size distributions which are difficult to characterise • non-spherical particles • difficulty with filtering and sample contamination ™ Better liquid to solid mass transfer and similar gas to liquid mass transfer. The liquid to solid mass transfer coefficient depends on the velocity of the solid relative to the liquid, known as slip velocity. The slip velocity is a strong function of particle size and relative density of the liquid. The swollen hydrated particle density of Amberlyst 15 has been reported as 0.927 g.mL-1 Leung et al (1986). Under various states of hydration the particle density will therefore be very close to that of the liquid. Hence, with small particles the danger exists that particles will be simply swept along by the liquid phase and no shear stress between particles and fluid will result, Fogler (1999). The use of a basket reactor at high Ω (rpm) results in high slip velocities between particles and fluid. ™ Samples taken from the basket reactor do not need to be filtered which means: • there is less chance of product loss in handling, which is important as small concentrations were to be analysed for, ≈ 0.001 mol.L-1.

Chapter 4: Kinetics of Hydration.... 120 • samples can be taken directly for the determination of isobutylene which is typically lost upon any form of additional handling due to its high volatility and low solubility.

Basket reactors have been successfully employed to investigate three phase systems of desulfurisation of dibenzothiophene, Myers and Robinson (1978); oxidation of , Njiribeako, Silveston and Hudgins (1978); and oxidation of sulfur dioxide, Pavko, Misic and Levec (1981). Most relevantly, Delion, Torck and Hellin (1986) employed a basket reactor for the determination of isobutylene hydration equilibrium constants in the presence of various solvents. Delion et al (1987) also investigated the kinetics of isopentene hydration within this basket reactor. They used a catalyst loading of 20 g.L-1 and obtained transport free kinetics. The main disadvantage of basket reactors is that the catalyst size cannot be investigated beyond a certain minium due to loss of catalyst from the baskets.

Chapter 4: Kinetics of Hydration.... 121 1/4" NPT 5

140 180

sintered ss plate

70 30

70 (a) (b) 100

M5 Allan

M5 bolt

d 30

20 25 (c)

mesh 15

Figure 4.01 Schematic of (a) the stirred basket slurry reactor employed, showing position of basket assembly and baffles (b) baffles and (c) baskets which the hold catalyst.

Chapter 4: Kinetics of Hydration.... 122 4.2-2 Catalyst Pretreatment Amberlyst 15 is supplied containing small amounts of byproducts resultant from the manufacturing process, which are typically removed using a pretreatment procedure before catalyst application. The Amberlyst 15 pretreatment procedure was based upon that developed by Leung et al (1986), it had been previously applied and validated by Zhang, Adesina, and Wainwright (UNSW (2003)) and involved the following steps: 1. The raw catalyst was sieved for a rough size range. 2. A 0.5 L sample was dried at 373 K for a period of 8 hr. 3. The catalyst was washed with 4, 1 L volumes of distilled water at 333 K. The colour of the wash water changed to light brown. 4. In an additional 1L of water the catalyst was heated and brought to 373 K under continuous mechanical stirring. It was held at this temperature for 0.5 hr. 5. The liquid was then decanted and the catalyst was allowed to cool. 6. It was then immersed in 1L of 10 wt % sulfuric acid solution in which it was held for 2 hr under continuous mechanical stirring. 7. The acid solution was then decanted and the catalyst washed with distilled water until the wash solution was found to be neutral. APS Litmus paper was used during the washing procedure to track the pH of the wash solution. Approximately 30, 1L volumes were used. 8. The catalyst was dried at 373 K for a period of 8 hr and allowed to cool at ambient temperature and humidity. 9. It was sieved a second time for the desired particle size range.

The hydrogen ion exchange capacity of the raw and treated Amberlyst 15 were evaluated following a direct titration procedure as developed by Fisher and Kunin (1955). The solids content was determined by drying a sample of catalyst overnight at 383 K. The results of the analysis are summarised in Table 4.01.

Table 4.01 Solids Content and Activity of raw and pretreated Amberlyst 15. Amberlyst 15 % Solids Dry mmol eq H+.g-1 Raw, As Received 73 5.4 Pretreated 70 5.3

Chapter 4: Kinetics of Hydration.... 123 The solids content of the catalyst was found to decrease upon treatment possibly due to the removal of any material left over from manufacture. The activity of the resin before and after pretreatment was comparable with little change in hydrogen capacity. The measured capacity was also comparable to manufacturer’s value, 4.7 mmol eqH+.g-1 (Rohm and Haas), and that of Boz et al (2004) of 5.1 mmol eqH+.g-1, using the same method.

4.2-3 Analytical Procedure TBA, MET, and isobutylene (list of materials given in Appendix A1) were analysed by gas chromatography (GC) using a Shimadzu GC-8A equipped with IFD detector. An Alltech manufactured and recommended HaySep-Q column was employed with particle size 60/80 mesh, column diameter 1/8”, length 8’ and Tmax 523 K. The GC and integrator settings are summarised in Table 4.02. A liquid sample of 2 μm was delivered using a gas tight micro-syringe. was employed as carrier gas and method of external calibration was adopted. Each overall calibration was performed as a ratio of the unknown area, A, and that of the average analysis run standard, AS.

Table 4.02 GC and Integrator Settings GC-8A Settings Integrator Settings Variable Value Variable Value Primary 600 kPa Attenuation 3 Nitrogen/Carrier 170 kPa Stoptime 40 min Pressure Hydrogen Pressure 60 kPa Slope 150 Air Pressure 40 kPa Width 5 Inj/Det 483 K Min. Area 0 Temperature Column 513 K Temperature Range 102 Attenuation 1 Polarity Plus

The presence of the solvent EG complicated the analysis by creating a strong non- linearity at low concentrations of TBA, by elevating the effective baseline. Thus, the calibrations were performed in the matrix at 30, 60 and 90 mol% EG. The EG also extended the sample time to 40 min. The correlations developed for TBA analysis were presented in Table 4.03.

Chapter 4: Kinetics of Hydration.... 124

Table 4.03 Calibrations for Analysis of TBA in Water/EG by FID Analysis Calibration R2 Eq. No.

Range: 0.0002 < CTBA < 0.2 TBA in 0.9981 (4.26) ⎛⎞ATBA Water CTBA = 0.1196⎜⎟ ⎝⎠AS

TBA in 2 0.9906 (4.27) ⎛⎞AA ⎛⎞ TBA TBA EG 30 CTBA =+0.1286⎜⎟ 0.0346 ⎜⎟ ⎝⎠AASS ⎝⎠ mol % TBA in 2 0.9848 (4.28) ⎛⎞AA ⎛⎞ TBA TBA EG 60 CTBA =+0.1803⎜⎟ 0.0122 ⎜⎟ ⎝⎠AASS ⎝⎠ mol % TBA in 2 0.9902 (4.29) ⎛⎞AA ⎛⎞ TBA TBA EG 90 CTBA =+0.0931⎜⎟ 0.0038 ⎜⎟ ⎝⎠AASS ⎝⎠ mol %

Overall 2 na (4.30) ⎛⎞AATBA ⎛⎞TBA CaTBA =+12⎜⎟ a ⎜⎟ ⎝⎠AASS ⎝⎠ 2 axx1 =−0.5903EGE + 0.6384 G

ax2 =−0.0514EG + 0.0477

The MET peak had a longer elution time than that of EG peak and was not affected in the manner of TBA peaks, thus a simple linear correlation sufficed at all EG concentrations, Table 404.

Table 4.04 Calibrations for Analysis of Ethylene Glycol mono-tert-Butyl Ether Analysis Calibration R2 Eq. No. MET in 0.9961 (4.31) ⎛⎞AMET W/EG CMET = 0.1391⎜⎟ ⎝⎠AS

For the analysis of isobutylene its peak areas were calibrated against literature on solubilities in water reported by Kazanskii et al (1959). These reference solubilities had been determined gravimetrically in a bomb type vessel under varying pressures of isobutylene and temperature. Experiments were conducted in a temperature controlled slurry reactor to gain the variation of isobutylene solubility with normalised GC peak area, Table 4.05. Leung et al (1987), and Zhang, Adesina and Wainwright Chapter 4: Kinetics of Hydration.... 125 (UNSW(2002)) had previously applied this method. The benefit of the method is that it allows for more complex mixtures and higher temperatures to be used than are possible in the gravimetric method alone.

Table 4.05 Calibration for Dissolved Isobutylene Analysis by FID Analysis Calibration R2 Eq No. IB in 0.4485 na (4.32) ⎛⎞AIB W / EG / TBA CIB = 0.0128⎜⎟ ⎝⎠AS

4.2-4 Experimental Procedure To allow for ease of operation and moderate operational cost the reactor was run batch in liquid and continuous in gas, Figure 4.02. The reaction rate was followed through the analysis of TBA, Section 4.2-3. Liquid samples of 0.5 mL were taken every 5 minutes for 40 minutes after experimental run initiation. Thus 8 samples were taken with a total volume of sampling of 4 mL, given that the initial charge was 350 mL this represented a 1.1 % change in the volume due to sampling, which was assumed to be negligible. Samples were refrigerated and analysed 1 hour after run completion. The reactor was allowed to cool and was opened up and cleaned after each run.

The variables of interest were temperature, liquid initial composition and gas phase composition. The temperature of each run was kept constant with the use of heating tape and temperature controllers. The liquid composition was varied in terms of initial reactor charge mixture in terms of water and ethylene glycol. The gas phase composition was varied through the use of mass flow controllers and a mixer. Although varying the solvent concentration inevitably varies gas solubility and hence isobutylene liquid phase concentration, it was decided to vary isobutylene partial pressure as well, so as to avoid coupling and interaction during regression and model determination. The mixtures for solvent runs of water and ethylene were prepared in advance in bulk (6 L) and kept in a sealed container.

Temperatures 341, 351 and 361 K were employed for this study. As shown in Figure 4.02 externally mounted heating tape was used to heat the reactor and maintain isothermal reaction conditions. The control method used was that of a separate

Chapter 4: Kinetics of Hydration.... 126 indicator/controller for the heating tape, with thermocouple secured between the heating tape and external wall of the reactor. An additional indicator was used to monitor liquid phase temperature. It was found that for ambient temperatures in the range 285 to 301 K, the heating tape control temperature needed to be set approximately 15 K above the desired reactor liquid charge temperature. The temperature was initially ramped up to 5 K below that desired using a setpoint of 383 K.

The experiments were to be performed in semi-batch mode with the gas phase continuous and the liquid phase batch. As such a condenser was used to reflux the vaporised product back to the reactor. The condenser was of jacket and coil-type and supplied with water at 293 K at a rate of 2 to 3 L.min-1.

The reaction runs were initiated with the addition of isobutylene to the gas feed stream. The liquid charge and catalyst were brought to the desired temperature using a flow of 0.2 L.min-1 of nitrogen. The nitrogen flow was adjusted down and the isobutylene flow was set to the desired amount. This point was taken as t = 0 s and the timing was initiated. This method was seen adopted because: • Isobutylene saturated the hot liquid rapidly. Based upon rapid stabilisation observed under the solubility runs and the general low level of solubility it was concluded that the kinetics of solution would not affect the reaction rates. Jayadeokar and Sharma (1993) had successfully employed this method for the determination of isobutylene etherification with ethylene and propylene glycol kinetics. • It minimised isobutylene wastage. • It allowed for optimum temperature control as the addition of catalyst results in temperature loss. • There was little scope for further reactor modification to include a catalyst addition device. • External catalyst addition could lead to release of vapours and catalyst loss due to sticking to wet surface and spillage.

The catalyst Amberlyst 15 had been pre-treated according to Section 4.2-2 and kept in a sealed container. It had been dry sieved to the particle size ranges described. The catalyst loading for each run was approximately 3 g (1.5 g in each basket), for a liquid

Chapter 4: Kinetics of Hydration.... 127 volume of 0.35 L giving 8.57 g.L-1. Fresh catalyst was used for each run. The catalyst size used had an average diameter of 450 μm (dry basis). A smaller size than this was not used to avoid excessive pressure drop.

Isobutylene solubility runs were conducted in much the same manner as the catalytic runs to determine rate of TBA formation. However, the catalyst was replaced with glass beads and the concentration of isobutylene was determined by GC analysis. The objective was to determine saturated steady state concentration.

For water free runs where the reaction rate was determined for the formation of mono- tert-butyl ethylene glycol ether (MET) by the reaction of isobutylene with ethylene glycol water free ethylene glycol and Amberlyst 15 were prepared. The ethylene glycol was brought to boiling point in order to eliminate water and charged in a well dried reactor. The weighed out catalyst was dried at 363 K overnight. The formation of MET was followed by GC analysis according to Section 2.2-3.

Chapter 4: Kinetics of Hydration.... 128

[01]

RPM [04]

[05] [02] [03] T1

T1 [08] [06] T2 [07] T2 [09] T

[18] [12]

PP [15] [10] [17]

[13]

P [16] [11]

[14] % Ch: 1 2 3 4

Figure 4.02 Stirred Basket Reactor Rig for Kinetics Determination

Key to Figure 4.02 Stirred Basket Reactor Rig for Kinetics Determination [01] Exhaust Fan of low suction. [02] Cooling Water Rotameter, TFM, Serial No. 9709006, water 0 to 10 L/min. [03] Condenser, Crown Scientific, QuickFit, 19/26 fittings, Jacketed Coil, CX6/22/SC, 4×10-2 m2 [04] Digital Stirrer, Heidolph, RZR-2021, 0 to 2000 rpm.

Chapter 4: Kinetics of Hydration.... 129 [05] Flexible Drive Connection, inhouse. [06] Reactor, SS316, inhouse, (Fig 4.01 detail). [07] Heating Tape, Isopad, Type TeMS6, 328W [08] Temperature Indicator, Shinko MCR controller, inhouse, fitted with K-type 1/16” ss thermocouple. [09] Temperature Controller, Shinko MCR controller, inhouse, fitted with K-type 1/16” ss thermocouple. [10] Nitrogen Supply Cylinder. [11] Isobutylene Supply Cylinder.

[12] Mass Flow Controller, Nitrogen Gas, Brooks 5850E, 1 L/min N2. [13] Mass Flow Controller, Isobutylene Gas, Brooks 5850E, 5 L/min Air. [14] Mass Flow Controller Control Box, Brooks Instrument 5898. [15] Shut Off Valve, Nitrogen Gas, Whitey 1/8”. [16] Shut Off Valve, Isobutylene Gas, Whitey 1/8”. [17] Gas Mixer, three port, 100 mL, brass. [18] Three Way Valve, Mixed Gas, Whitey 1/8”.

Operating Procedure for Kinetics Experiments The following procedure was employed for the determination of a single reaction rate point: 1. Weigh out using digital balance previously prepared Amberlyst 15 catalyst into the baskets of the reactor, aim at 1.5 g per basket and write down the actual mass to four places. 2. Attach baskets to basket holder mounted on the reactor shaft, which is connected to the reactor lid. 3. Prepare liquid charge by measuring out 0.35 L of the desired composition mix previously prepared and analysed. 4. Turn on nitrogen at cylinder [10], open shut off valve [15], set nitrogen flow at control box [14] to a setpoint of 5%. 5. Pour liquid charge into reactor [6] and assemble the reactor by replacing its lid and securing four holding nuts with Allen key. 6. Connect the flexible drive [5] to the digital stirrer [4]. 7. Turn on Exhaust Fan [1], at wall mounted switch.

Chapter 4: Kinetics of Hydration.... 130 8. Turn on Cooling water at Cooling Water Rotameter [2], set to 2 L.min-1. 9. Turn on digital stirrer and adjust the setpoint to Ω = 1400 rpm. 10. Ensure the two thermocouples are in place and secure. 11. Adjust the nitrogen flow to its experimental setpoint. 12. Set temperature controller [9], to desired setpoint and monitor during heating to desired liquid setpoint. 13. Take the t = 0 min sample using a 5 mL syringe and drawing a 0.5 mL sample after three wash draws and returns. 14. Prepare stopwatch, open isobutylene cylinder and shut off valve, [11] & [17]. 15. Set the desired isobutylene flow by switching the observed channel on the control box to isobutylene and adjusting the scaled knob till the desired setpoint is displayed. Start the timer at this point, t = 0 min. 16. Take samples every 5 min for 40 min. Use the three wash draw/return method each time. Place the samples in vials and refrigerate immediately. 17. Upon completion of sampling, shut down the isobutylene, turn off the heating tape, turn off the stirrer. Leave the cooling water running and allow the reactor to cool to room temperature before opening, discharging fluid and clean. 18. The regular cleaning procedure involves removing the baffles, which were washed separately. Filling the reactor with water at least three times and then once with boiling water. Wiping down the inside of the reactor with tissue paper. Once every set of experiments the bottom of the reactor can also be disassembled and the gas chamber cleaned and the sintered distributor retaped with teflon high density natural gas plumbing tape. 19. The lid is washed separately, the baskets emptied of catalyst, washed and oven dried at 393 K to remove moisture. 20. After cleaning the nitrogen can be turned off. 21. The liquid samples are now analysed for TBA and MET, Section 4.2-3.

Chapter 4: Kinetics of Hydration.... 131 4.2-5 Choice of Variables and Experimental Design The main aim of the study was to determine the kinetics of hydration in the presence of the proposed solvent ethylene glycol, as the hydration of isobutylene itself has been studied over the catalyst of interest. Of particular interest was the evaluation of the relative rates of reaction of water and ethylene glycol with isobutylene.

The solubility runs were conducted at 315, 335 and 355 K. These temperatures were dictated by the necessity of comparison to literature values given at these temperatures.

The solubility within pure and binary mixtures was evaluated at yIB 0.5, 0.75 and 1, while ternary mixtures were only evaluated at yIB = 0.5.

Transport runs were conducted to eliminate external mass transfer effects and characterise internal mass transport over the moderate particle size of 450 μm employed with the basket reactor. External mass transfer effects were evaluated in a system high in ethylene glycol as it was considered that basket penetration and gas to liquid mass transfer would be poorest in the high viscosity, high density glycol. Further, the lowest temperature of the study was 341 K. It was assumed that should external mass transfer limitations be avoided for the poorest conditions then other conditions should also be free of mass transport artefacts. The internal mass transfer runs were conducted at high concentrations of water as the catalyst preferentially adsorbs water and all runs would be conducted with high water content.

The experimental design used was essentially that of complete matrix variation over the ranges of interest, Table 4.06. In light of the aims, the following experimental approach was devised: • Solubility Runs. To determine the solubility of isobutylene in water, ethylene glycol and TBA and their binary and ternary mixture. • Transport Runs. Varying rate of stirring to eliminate external mass transfer. Examine the effect of three different particle sizes to gain an estimate of the effectiveness factor. • Water runs. Measuring the reaction rate in the absence of solvent establishing the base case for evaluation of the effect of solvent and allowing for comparison with previous work in the field.

Chapter 4: Kinetics of Hydration.... 132 • Solvent runs. Measuring the effect of solvent concentration upon reaction rates and

selectivity. These included a series across the full solvent range at xEG = 0.1

resolution at 361 K and yIB = 0.75.

Table 4.06 Kinetic Study Experimental Range and Design Manipulated Variables Setpoints Solubility Runs Isobutylene mole fraction yIB 0.50, 0.75, 1.00 Temperature (K) 315, 335, 355 binary xi 0.5 ternary xi 0.17, 0.33, 0.67 Transport Runs Ω (rpm) 600, 800, 1000, 1200, 1400, 1600

dp (μm) 350, 450, 550 Water Runs Temperature (K) 341, 351, 361 Isobutylene mole fraction yIB 0.25, 0.50, 0.75 Solvent Runs Temperature (K) 341, 351, 361 Isobutylene mole fraction yIB 0.25, 0.50, 0.75 Ethylene glycol xEG 0.1 to 1, bulk: 0.3, 0.6, 0.9

4.2-6 Calculation Procedure In isobutylene solubility runs, concentrations were calculated from the raw GC areas as they were obtained and plotted. Steady state was assumed once successive GC areas were within 10 % of each other. Typically, the average of the last three to four measurements were used as the reported solubility.

The reaction rate of each run was determined by following the appearance of the product TBA. The observed rates of reaction were found to be constant over the shorter runs of small amounts of change in reagent concentration (1.0 × 10-03 mol of water converted). Thus, the rate was determined using linear regression. The initial product concentration was in all cases 0 mol.L-1, with no occurrences of contamination and nonlinear low concentration effects having been accounted for in the GC calibrations. Thus, the regression used the origin as a data point. The slope of the fitted line gave the reaction rate based upon the liquid volume, rV:

Chapter 4: Kinetics of Hydration.... 133 CrtTBA= V (4.33) which was used to calculate the rate based upon mass catalyst by the following equation at a given temperature: η()rV rmolgs ()..−−11 = VL (4.34) m where: η effectiveness factor ( - ) -1 -1 rv reaction rate based on liquid volume (mol.L .s )

VL liquid volume, (L) m catalyst mass, (g)

While the reaction rates were reported in mol.g-1.s-1, it is noted that the measured equivalent acidity of the pre-treated catalyst used was 5.3 × 10-03 mol H+/g, should comparison against another catalyst be required.

4.3 Results and Discussion In the following section the results obtained were presented and discussed in terms of the effects of the manipulated variables, comparison to literature values and correlations. Isobutylene solubility is discussed first. Existing correlations were applied and evaluated for fit and appropriateness. These results were used to interpret the reaction rate of isobutylene hydration to TBA. The was carried out in the presence of ethylene glycol and its effect upon reaction rate and side reaction to MET are discussed. Finally, kinetic models are proposed for the hydration of isobutylene to TBA in the presence of the solvent ethylene glycol on the basis of literature and general observations and are evaluated for statistic and thermodynamic consistency. Tables of the results obtained using the experimental design of Section 4.2- 5 are presented in Appendix A2.

4.3-1 Isobutylene Solubility The solubility of isobutylene within mixtures of TBA/W/EG was measured in dedicated runs, independently of reaction runs, in order to improve the accuracy of measurement.

Chapter 4: Kinetics of Hydration.... 134 The correlations gained were used to allow for kinetic modelling based upon bulk liquid phase concentrations of the relevant components including isobutylene. Due to the amphiphilic nature of TBA in possessing both alkyl and hydroxyl groups in a tight arrangement that allow for good miscibility with polar solvents such as water and ethylene glycol strong nonlinearities were observed in isobutylene solubility. These features being of particular interest were also accounted for in the correlations developed.

Solubility expressed as molar fraction of isobutylene as a function of isobutylene partial pressure is reported in Figure 4.03. Generally, solubility increased with non-polarity of the liquid phase and isobutylene gaseous molar fraction and decreased with temperature as was anticipated. It was observed that the measured isobutylene liquid phase molar fraction, xIB, varied linearly with isobutylene gas phase partial pressure, yIBP. Thus, Henry’s law appears applicable and the slope of the curve at each temperature is the Henry’s law constant corresponding to:

yAAPHx= (4.35)

Chapter 4: Kinetics of Hydration.... 135

1.2 1.0 0.8 0.6 316 K P (atm)

IB 0.4 334 K y 0.2 353 K 0.0 0.0E+00 2.0E-05 4.0E-05 6.0E-05 8.0E-05

xIB (a) water

1.2 1.0 0.8

0.6 316 K P (atm)

IB 0.4 334 K y 0.2 353 K 0.0 0.0E+00 2.0E-04 4.0E-04 6.0E-04 (b) EG xIB

1.2 1 0.8 0.6 316 K P (atm) 0.4 IB

y 334 K 0.2 353 K 0 0.0E+00 2.0E-03 4.0E-03 6.0E-03

xIB (c) TBA

Figure 4.03 The solubility of isobutylene, xIB, versus partial pressure yIBP and temperature for (a) water, (b) ethylene glycol and (c)TBA.

Chapter 4: Kinetics of Hydration.... 136 The relationship of the Hi values obtained with temperature could be determined from

Figure 4.04 which suggested Arrhenius type variation in Hi with T and thus for the pure components H were correlated by:

HTW =−4607831exp( 1801/ ) (4.36)

HTEG =−32249exp( 869 / ) (4.37)

HTTBA =−7257exp( 1197 / ) (4.38)

The adequate fit of these correlations is demonstrated in Figure 4.04.

For pure components using the approximation of Henry’s law for the prediction of the behaviour of gas solubility, Prausnitz (1969) showed that:

ΔΔshoo⎛⎞1 ln xeq =−⎜⎟ (4.39) R RT⎝⎠

where xeq is the isobutylene mole fraction in the liquid in equilibrium with isobutylene gas at 101.3 kPa. Thus, a plot of ln xeq versus 1/T should give a straight line where slope and intercept provide values of the standard partial molar enthalpy of solution, Δho, and the standard partial molar entropy of solution, Δso. The results of applying Equation 4.39 are shown in Figure 4.05. The agreement between data from the present study and those of Leung et al is evident in Figure 4.05. The values of Δho and Δso obtained are reported in Table 4.06. It can be seen that the absorption of isobutylene is an exothermic process, consistent with the observed decrease of solubility with increased temperature. As can be seen from Table 4.07 the addition of TBA gives rise to increasingly positive enthalpies of solution.

Chapter 4: Kinetics of Hydration.... 137 3.0E+04 Water 2.5E+04 EG TBA 2.0E+04

1.5E+04 H (atm) 1.0E+04

5.0E+03

0.0E+00 0.0028 0.0029 0.003 0.0031 0.0032 1/T (1/K)

Figure 4.04 The effect of temperature on the Henry’s constant obtained, showing an Arrhenius type behaviour.

-4 W Leung (1987) -5 W -6 EG )

IB -7 TBA -8 ln ( x W/EG 1:1 -9 W/TBA 1:1 -10 -11 EG/TBA 1:1 0.0028 0.003 0.0032 0.0034 1/T (1/K)

Figure 4.05 Determination Δho and Δso of isobutylene absorption in W/EG/TBA. Also plotted are the data of Leung et al (1987) for comparison.

Chapter 4: Kinetics of Hydration.... 138

Table 4.07 Enthalpy and Entropy of Isobutylene Solubility o o Solvent Δh Δs solution Δh mixing Cohesive solution J.mol-1 J.mol-1 Energy Density** J.mol-1 @ 298.15 K J.mol.mL-1 W -15615 -130 4935 2095.93 EG -8970 -92 11580 857.86 TBA -7947 -69 12603 389.69 W.EG 1:1 -5788 -96 14762 W.TBA 1:1 -3815 -69 16735 EG.TBA 1:1 -5757 -69 14793 **Gu et al (2004)

The partial molar enthalpy change for solution results from the addition of the enthalpy of condensation, Δhc and the molar enthalpy of mixing Δhmix:

o Δ=Δ+Δhhhcmix (4.40)

The enthalpy of condensation is practically constant with a value of –20 550 J.mol-1 Leung et al (1987). Thus, the increase in Δho, which results from the addition of TBA, stems from an increase in Δhmix. According to Leung et al (1987a) the effect of the alcohol in a mixture with water is to increase the cohesive energy density of the liquid solution relative to that of pure liquid isobutylene. The endothermic nature of mixing in the pure components arises as more energy is required to break the molecular attraction between water-water, ethylene glycol- ethylene glycol, and TBA-TBA than is made available by the incorporation of isobutylene within these molecular networks. TBA forms highly cohesive solutions with W and EG due to its tight cruciform structure, which promotes a high degree of molecule intermeshing and its centrally positioned – OH group, which forms a continuum in hydrogen bonded networks.

Reported in Table 4.07 are the cohesive energy densities of the pure solvents at 298.15 K which are used as a standard measure of solvent-solvent interaction, Gu et al (2004). The cohesive energy density (ce) is defined:

Chapter 4: Kinetics of Hydration.... 139 (Δ−hRTvap ) ce = (4.41) V

where Δhvap is the enthalpy of vaporisation and V is the molar volume of the solvent. It increases with molecular interaction such as hydrogen bonding and influences bulk properties such as viscosity and surface tension. Both enthalpy and entropy of isobutylene solution in the pure solvents varied linearly with the cohesive energy densities of the pure solvents. The change in entropy of isobutylene solution was negative, which is consistent with condensation. The change was most negative in magnitude when isobutylene dissolved in the highly polar and highly associated structure of water.

Figure 4.06 demonstrates the strong positive enhancement effect of TBA upon solubility of isobutylene in mixture of TBA and water and TBA and ethylene glycol. Figure 4.06 (a) showed that solubilities within water and ethylene glycol were additive as the solubility points within the 1:1 mixture fell half way between the solubilities within the pure components. Conversely, Figure 4.06 (b) and (c) show that the solubilities within mixtures containing TBA are highly skewed towards TBA. This suggests that isobutylene associates strongly with TBA in solution. This association depends upon existing structures within the solution and the possibility and freedom of new structure formation.

Water miscible alcohols such as methanol and ethanol are known to exhibit microscopic structure formation such as clustering and aggregation in aqueous solutions, Yoshida et al (2002). These have been shown to have pronounced effects upon the thermal unfolding of nucleic acids and proteins and the action of surfactants, Freda et al (2001). Amongst mono-hydroxyl alcohols that are miscible with water in any proportion, TBA posses the biggest alkyl groups and hence exhibits the largest hydrophobic effect. The structure of TBA-water mixtures was determined using neutron diffraction techniques by Bowron et al (1998) and infrared techniques by Freda et al (2001). The findings of these experimental bond stretching investigations have also been modelled using statistical thermodynamics tools and in particular the reference interaction site model with the closure approximation by Yoshida et al (2002). Their combined efforts indicate

Chapter 4: Kinetics of Hydration.... 140 that the most probable structure, that with the highest frequency of occurrence, of TBA- water solutions changes gradually as their proportions are varied. In dilute TBA solutions the hydroxyl group of a TBA molecule is included into the water tetrahedral like hydrogen bonding network forming a cage around the TBA hydrophobic tert-butyl group. With the rise of the TBA mole fraction in aqueous solutions, (TBA > 0.2), TBA molecules progressively cluster by the tert-butyl groups due to their hydrophobic attraction in water, while the TBA hydroxyl groups remain included in the hydrogen cage of water around the TBA aggregates. In concentrated TBA solutions (TBA > 0.4) zigzag hydrogen bonded chains arise which incorporate the water molecules present. According to Yoshida et al (2002) these structures are continuously breaking and reforming and exchanging molecules with thermal fluctuations.

The formation of structures of TBA with water provide for clusters of alkyl groups, which promote the solution of isobutylene even at low TBA concentrations and may occupy the polar hydroxy groups of water and or ethylene glycol effectively shielding isobutylene from their influence. This may explain the nonlinear solubility in the presence of TBA observed in Figure 4.06. It can be suggested that the aggregate type structures formed according to Yoshida et al (2002), dissolve isobutylene with an affect akin to micelle emulsification of highly non-polar fats in aqueous solutions. It can be suggested on the basis of the similar behaviour of isobutylene solubility in ethylene glycol/TBA mixtures that similar structures may form in this system to those present in the water/TBA system as previously discussed.

For a mutlicomponent system O’Connell and Prausnitz (1964) proposed:

nnn ss ln HP2,mix () 1 =−∑∑∑ xHPjln 2, j () 1 α jk xx j k (4.42) jj==11k>1 jj≠≠22k≠2 where component 2 is the solute and component 1 is the reference solvent (used to reference the pressure) and xj and xk are molar fractions of the components of the solution. The deviation from combination of Henry’s constants by mole fraction is represented by interaction parameters in the form of Margule’s constants, α, which are

Chapter 4: Kinetics of Hydration.... 141 typically evaluated through the direct regression of experimental data. Applying the 2- way interaction model of Equation (4.42) to the present data the following set of equations was obtained:

lnHxHxHxH=+ ln ln +ln + IB... mix W IB W EG IB EG TBA TBA (4.43) ααW.. EGxx W EG++ W TBA xx W TBA α EG . TBA x EG x TBA where:

−03 αWEG. =×3.026 10T − 14.14 (4.44)

−03 αWTBA. =−7.281 × 10T − 14.88 (4.45)

−02 α EG. TBA =−2.069 × 10T + 3.53 (4.46) with a correlation coefficient R2 = 0.937. However if a three way interaction was considered,

lnHxHxHxH=+ ln ln +ln + IB... mix W IB W EG IB EG TBA TBA (4.47) ααW..... EGx Wxxxxxxxx EG++ W TBA W TBA α EG TBA EG TBA + α W EG TBA W EG TBA where:

−03 αWEG. =×4.818 10T − 14.154 (4.48)

−02 αWTBA. =−3.321 × 10T + 0.270 (4.49)

−02 α EG. TBA =−1.889 × 10T + 3.514 (4.50)

−03 αWEGTBA.. =−5.489 × 10T − 14.89 (4.51) with a slightly improved coefficient R2 = 0.941. As a result a 2-way interaction model would be sufficient in subsequent interpretation of the kinetic data.

Chapter 4: Kinetics of Hydration.... 142 6.E-03

) 5.E-03 W -1 4.E-03 EG 3.E-03 xEG = 0.5 (mol.L 2.E-03 IB

C 1.E-03 0.E+00 310 320 330 340 350 360 T (K) (a)

0.05 W

) 0.04 TBA -1 0.03 xTBA = 0.5

(mol.L 0.02 IB

C 0.01 0 300 320 340 360 T (K) (b)

0.05 EG

) 0.04 TBA -1 0.03 xTBA = 0.5

(mol.L 0.02 IB

C 0.01 0 300 320 340 360 (c) T (K)

Figure 4.06 Binary mixtures of water, EG and TBA: (a)W/EG, (b) W/TBA and (c) EG/TBA, demonstrating the strong positive enhancement effect of TBA upon isobutylene solubility expressed here in mol.L-1.

Chapter 4: Kinetics of Hydration.... 143 4.3-2 Transport Analysis Transport steps comprise of those external and internal to the catalyst particle. The reactor chosen, rotating basket reactor, allowed for the external mass transfer limitations to be eliminated. However, catalyst size constraints required the internal mass transfer to be characterised. Internal effects were compensated for using an effectiveness factor.

In order to ensure that external mass transfer was not rate limiting a method of elimination and operation within a mass transfer limitation free regime was applied. The speed of revolution, Ω, of the basket impeller packed with catalyst of dp = 450 μm, was varied in the range 600 to 1600. It is evident from Figure 4.07 that further increase of Ω beyond 1000 rpm has little influence upon the observed reaction rate. As such further experiments were conducted at a conservative, yet comfortable, value of Ω = 1400 rpm.

12.000

10.000 ) -1 .s

-1 8.000

(mol.g 6.000 +09

.10 4.000 TBA r 2.000

0.000 0 500 1000 1500 2000 Ω (rpm)

Figure 4.07 The elimination of the effect of external mass transport artefacts upon -1 -1 observed reaction rate, rTBA (mol.g .s ), through adequate stirring speed, Ω (rpm).

Chapter 4: Kinetics of Hydration.... 144 In order to ensure that internal mass transfer was not rate limiting a method of characterisation and evaluation was applied. Previous studies have indicated that pore- diffusion resistance necessarily accompanies isobutylene hydration due to swelling of catalyst particles in the aqueous environment. As a result, even with small particles (on a dry basis), it is still important to characterise the basket reactor for internal mass transfer effects. Several investigators (Gupta and Douglas, Ihm et al, and Velo et al) have reported that the effective diffusivity of isobutylene decreases with temperature. Ihm et al suggested that the Amberlyst 15 particles consist of gel microparticles and a macroporous space. They concluded that the internal transport resistance is contributed primarily by the diffusion in the macroporous region since the effectiveness factor in this space is essentially the same as the experimentally measured value while that for the gel particles is practically unity.

The reaction rate of isobutylene hydration, following the production of TBA, was determined for particles of size 350, 450 and 550 μm at 361 K. The runs were conducted in duplicate and the results reported in Figure 4.08.

2.0E-07 )

-1 1.5E-07 .s -1 1.0E-07 (mol.g

TBA 5.0E-08 r

0.0E+00 300 400 500 600

dp (μm)

Figure 4.08 The determination of internal mass transfer parameters for wet

Amberlyst 15of dp: 350, 450 and 550 μm at 361 K.

Chapter 4: Kinetics of Hydration.... 145 These runs were analysed using the Weisz-Prater criterion given as:

2 2 −rRAcρ ηφ1 = (4.52) DCeAS

A first order reaction rate in isobutylene was assumed based upon available literature and was confirmed in later runs. The internal effectiveness factor for a first-order reaction in a spherical catalyst pellet is given by:

11⎡ ⎤ ηφ=−⎢coth 3 ⎥ (4.11) φ ⎣ 3φ ⎦ thus combining Equations (4.52) and (4.11) the following ratio was obtained:

2 −−rR2 2φφ 12coth 12 1 2 = = A (4.53) −−rR1 1φφ 11coth 11 1 where the subscript refers to two experimental runs at different diameters. Substituting for φ11 = Bφ12 and rearranging gives:

( A −1) φ12 = (4.54) ()ABφ12coth Bφφ 12− 12 coth φ 12

which is in a form suitable for solution of φ12 by initial estimate and subsequent iteration. The results of the analysis are reported in Table 4.08. The estimate of a lower magnitude was adopted, hence with dp = 450 at T = 361 K the effectiveness factor was approximated as η = 0.39.

Table 4.08 Calculation of Internal Mass Transfer Parameters According to Equations (4.53) to (4.55)

dp (1) dp (2) A B φ11 φ12 η11 η12 350 450 1.3550 0.7778 5.1171 6.5791 0.4717 0.3867 450 550 1.2745 0.8182 5.2417 6.4066 0.4632 0.3952 350 550 1.7270 0.6364 4.6643 7.3296 0.5054 0.3535

Chapter 4: Kinetics of Hydration.... 146 The dependence of effectiveness factor upon dp was correlated with the following expression:

−−05 0.791 η =×1.825 10 d p (4.55)

The current data corroborates the findings of Leung et al (1986). Their effectiveness factor data for dp 450 when correlated with temperature by

−03 ⎛⎞1331.141 η =×9.76 10 .exp⎜⎟ (4.56) ⎝⎠T suggests a value of η = 0.389 at T = 361.

Recognising that effectiveness factor will need to be a function of temperature the dependence of the effectiveness factor upon temperature developed by Leung et al (1986) was adopted. For all available data of the current study and that of Leung et al (1986) the following correlation was proposed allowing for η to be estimated for

Amberlyst 15 for different T and dp:

−0.694 ⎛⎞1343.510 η = 0.652d p .exp⎜⎟ (4.57) ⎝⎠T

2 where the units of dp are μm. The correlation returned R = 0.986. Given that high amounts of water were present for all runs, xw ≥ 20 mol %, except for pure solvent runs, the catalyst should remain well hydrated resulting in wet catalyst and similar internal diffusivities. The concentration of TBA, which can affect internal mass transfer of isobutylene, remained low for all runs CTBA ≤ 0.005 mol/L. Thus, Equation (4.57) was used with confidence for the kinetic study.

For runs containing only the solvent ethylene glycol the work of Jayadeokar and Sharma (1993) is referenced. They found that for Amberlyst 15 for the reaction of isobutylene and ethylene glycol, for particles of size dp = 420 and 600 μm, operating at 323 - 343 K, internal mass transfer rate did not influence the observed reaction rate, thus η = 1.

Chapter 4: Kinetics of Hydration.... 147 4.3-3 Pure Water Runs

The solvent free kinetics, xEG = 0, were investigated to provide a bench mark for comparison with solvent-based reaction rates and comparison to previously published data. Figure 4.09 demonstrated the linear influence of isobutylene, yIB, upon reaction rate. The overall linearity was in agreement with previously published findings suggesting first order behaviour in isobutylene. The reaction rate was also found to increase exponentially with temperature.

1.2E-06 341 K 1.0E-06 351 K

) 8.0E-07 361 K -1 .s -1 6.0E-07

r (mol.g 4.0E-07

2.0E-07

0.0E+00 0 0.0005 0.001 0.0015 0.002 -1 CIB (mol.L )

Figure 4.09 The reaction rate of TBA formation with isobutylene liquid phase concentration and temperature.

The pseudo first order reaction rate constants, k’, were correlated by:

−−11 ⎛⎞−7195.4 rmolgsTBA (). .= 351178exp⎜⎟CIB (4.58) ⎝⎠T

The activation energy 59.82 kJ.mol-1, compared well with 61.5 to 73.7 kJ.mol-1 over Amberlyst 15 reported in the literature, Leung et al (1986) and Ihm et al (1988). Figure 4.09 shows that the reaction rate constants given by Equation (4.58) fall within the Chapter 4: Kinetics of Hydration.... 148 range of constants predicted by published kinetic models. The constants compare best to those of studies conducted at similar temperatures: Ihm et al. (1988) (323 to 353 K), and Zhang et al. (2003) (323 to 353 K). The model of Leung et al. (1986) (303 to 333 K), does not extrapolate as well with higher temperatures, likely due to differences in isobutylene solubility correlation with temperature, Figure 4.05. The use of the Arrhenius expression and the exponential behaviour of reaction rate in temperature results in very high sensitivity to temperature and hence to error in experimental measurement of reaction rate. In turn this comparison demonstrates the importance of acquiring kinetic data as close as possible to the temperature of interest for further use, eg. catalytic distillation.

0.0008

0.0007 Current Study 0.0006 Leung (1986)

) Ihm (1988)

-1 0.0005

.s Zhang (2003) -1 0.0004

0.0003 k' ( L.g

0.0002

0.0001

0.0000 0.0026 0.0028 0.003 0.0032 0.0034 1/T (K-1)

Figure 4.10 Comparison of pseudo first order rate constants obtained in the current study to published data at temperatures particular to each

Chapter 4: Kinetics of Hydration.... 149 4.3-4 Empirical Kinetic Model for Ethylene Glycol Mediated Isobutylene Hydration The selection of an appropriate solvent for an entrainer to break the TBA-water azeotrope is important in this catalytic distillation study. The solvent must be effective, non-toxic, relatively unreactive and easy to recover. Amongst other suitable solvents ethylene glycol was chosen because of its relatively lower reactivity and higher effectiveness, as discussed in Section 3.2. Its feasibility and functionality as a reaction solvent has yet to be characterised. Very few solvents will remain inert in this water, TBA, isobutylene, and Amberlyst 15 system. They contributed to the reactive groups of the unsaturated double bond of isobutylene, and the hydroxy groups of water and TBA as well as a general purpose strongly acidic catalyst Amberlyst 15. It was concluded that the most favourable solvent, ethylene glycol, should be trialed and the extent and nature of the resulting reaction network characterised. In favour of this approach were: ™ the high affinity of Amberlyst 15 for water, Aiouache and Goto (2003). ™ the possibility that the solvent selection strategy may be extended to include obviously reactive solvents, but that selectivity may be controlled through temperature, mixing and separation including the suppression of byproduct removal, formation of azeotropes and recycle.

The parallel series reaction network, which can be established through the use of ethylene glycol as a solvent, includes: 1. Hydration of isobutylene to TBA, Equation (4.59) 2. Reaction of isobutylene with ethylene glycol to give the mono-tert butyl ethylene glycol ether (MET), Equation (4.60) and 3. Further reaction of the MET with isobutylene to give the di-tert-butyl ethylene glycol ether (DET), Equation (4.61)

h+ CH C=+ CH H O⎯⎯→ CH COH (4.59) ()322323←⎯⎯ ()

h+ HO CH OH+= CH C CH⎯⎯→ HO CH OC CH (4.60) ()23222 () ←⎯⎯ ()() 23 23

h+ HO CH OC CH+= CH C CH⎯⎯→ CH CO CH OC CH (4.61) ()()()2332232←⎯⎯ ()()() 323 323

Chapter 4: Kinetics of Hydration.... 150 For the purpose of the current study only the first two reactions were considered. The reaction runs were kept short with concentration low such that DET would not be formed to any appreciable extent. Jayadeokar and Sharma (1993) observed a lag of 90 min before the appearance of DET, while operating at 333 K and using a substantial slurry loading of 10 wt % or 100 g.L-1. Additionally, forward reaction rate constants for MET and DET formation differed by an order of magnitude. On the basis of this information the network was reduced to two parallel reactions Equation (4.59 and 4.60).

The effects of ethylene glycol addition upon isobutylene hydration can be seen in Figure

4.11 for runs conducted at different solvent mole fractions (xEG: (a) 0.32, (b) 0.55 and

(c) 0.81); different yIB (0.25, 0.5 and 0.75) and T (341, 351 and 361 K). It is evident that as the solvent concentration is increased the reaction rate, rTBA, decreases even though there is some improvement in isobutylene solubility. The hydration remains first order for lower loading of ethylene glycol, however departure from first order can be noticed as xEG is increased. The effect is most noticeable at higher temperatures. It may simply result from the reduction of the degree of excess of water for these runs.

The rate of reaction of formation of MET in the presence of water for xEG = 0.81 also follows first order behaviour in isobutylene, Figure 4.12. The reaction rate increased with temperature and an activation energy of 63.13 kJ.mol-1, was obtained over hydrated Amberlyst 15. It is immediately evident from a comparison of Figures 4.11 (c) and 4.12 that the reaction rates of formation of the alcohol and the ether are different by almost an order of magnitude. This suggests that even over the hydrated catalyst the sites and catalytic action of Amberlyst 15 may be different for the two reactions.

In view of the solubility of isobutylene in both water and ethylene glycol, there is a coupling of the concentrations of these species in the determination of reaction kinetics.

To investigate this phenomenon, Figure 4.13 (a) plots the rate of TBA formation, rTBA, against the product CIBCW in the presence of ethylene glycol. The linear relationship:

rkCCTBA= ' IB W (4.62)

Chapter 4: Kinetics of Hydration.... 151 suggests that the kinetics of TBA formation is first order with respect to both isobutylene and water. However, Figure 4.13 (b) reveals that the rate of MET formation is a nonlinear function of the concentration product CIBCEG in the presence of water. This power law relation:

α α rkCCMET= IB EG (4.63) yields α ≈ 2 indicating that the reaction was possibly between associated molecules of isobutylene and ethylene glycol.

The studies of Velo et al (1988), Leung et al (1986) and Gupta and Douglas (1967) all propose first order kinetics in both isobutylene and water for the formation of TBA. The study of Jayadeokar and Sharma (1993) suggests first order behaviour in both isobutylene and ethylene glycol for formation of MET. It is evident from Figure 4.13 (a) that in the presence of ethylene glycol, the hydration of isobutylene remains almost linear with respect to the product CIBCW. Thus, the predominant effect of ethylene glycol upon the hydration of isobutylene is that of linear dilution. It dilutes the concentration of water and through adsorption dilutes the number of active sites in a competitive manner with water.

Conversely the glycolisation of isobutylene to MET in the presence of water, Figure

4.13 (b), is highly curved in CIBCEG. This curvature is likely to be the result of the strong retardation of the reaction of ethylene glycol with isobutylene by water. The last point of the plot (CIBCEG = 0.08) corresponds to water free MET formation. The reaction rate drops sharply with the addition of water in the direct of the origin.

The difference between pure ethylene glycol rate of formation of MET, and pure water rate of formation of TBA (last points of each plot Figure 13 (a) and (b)) suggests that the two may occur at different sites within Amberlyst 15. Further, with the retardation of the reaction of ethylene glycol with isobutylene it is likely that water affects the sites which best accommodate this reaction. Gates and Rodriguez (1973) in their study upon the dehydration of TBA pointed out that the nature of the acid sites and consequently their activity changes with the concentration of water. They demonstrated that upon the addition of water there is a transition from catalysis by bound –SO3H groups, to

Chapter 4: Kinetics of Hydration.... 152 catalysis by hydrated protons in the matrix. They showed that catalysis by the –SO3H groups is much more rapid than catalysis by hydronium ions. The first order reaction rate constant for the dehydration of TBA they found for water free catalysis was 40 times the first order constant over the hydrated catalyst. Thus, the inhibition of water upon the reaction of ethylene glycol and isobutylene is likely to be caused by a reduction in the concentration of ethylene glycol and an alteration of the sites to a form found to be less active. Hence the steep drop in reaction rate of MET formation from that of pure ethylene glycol and isobutylene with the addition of water.

Chapter 4: Kinetics of Hydration.... 153

8.0E-07 341 K 7.0E-07 351 K 6.0E-07 361 K )

-1 5.0E-07 .s -1 4.0E-07

3.0E-07 r (mol.g 2.0E-07

1.0E-07

0.0E+00 0 0.001 0.002 0.003 0.004

-1 CIB (mol.L ) (a)

6.0E-07 341 K 5.0E-07 351 K 361 K

) 4.0E-07 -1 .s -1 3.0E-07

r (mol.g 2.0E-07

1.0E-07

0.0E+00 0 0.001 0.002 0.003 0.004 0.005

-1 CIB (mol.L ) (b)

2.5E-07 341 K

2.0E-07 351 K 361 K ) -1

.s 1.5E-07 -1

1.0E-07 r (mol.g

5.0E-08

0.0E+00 0 0.001 0.002 0.003 0.004 0.005 0.006

-1 (c) CIB (mol.L )

Figure 4.11 The influence of ethylene glycol upon isobutylene hydration at (a) xEG

0.32 (b) xEG 0.55 and (c) xEG 0.81

Chapter 4: Kinetics of Hydration.... 154 In order to account for the concurrent influence of water, isobutylene and ethylene glycol during isobutylene hydration in a ternary system the data of all runs were fitted to a power law:

−−1 1 ⎛⎞−4091 0.12 rmolgsTBA ()..= 1.0exp⎜⎟CCCIB W EG (4.64) ⎝⎠T

The slight positive dependence on ethylene glycol concentration could be due to the enhanced solubility of isobutylene in the presence of ethylene glycol in the aqueous reaction mixture. Treatment of the rate of reaction of ethylene glycol and isobutylene to MET in a similar manner gave an adequate fit (R2 = 0.8955) and the following expression:

−−1 1 ⎛⎞−2964 −0.89 rmolgsMET (). .= 1.0exp⎜⎟CCCIBEGW (4.65) ⎝⎠T

The inhibition effect of water upon the formation of MET evident in Figure 4.13 was captured by the negative exponent in CW of Equation 4.65. In order to optimise reaction conditions for TBA, its selectivity ratio, SRTBA, was defined:

rTBA ⎛⎞−1128 1.89− 0.89 SRTBA==1.0exp⎜⎟ C W C EG (4.66) rTMET ⎝⎠ and alternatively:

0.89 ⎛⎞−1127 ⎛⎞CW SRTBA=1.0exp⎜⎟⎜⎟ C W (4.67) ⎝⎠TC⎝⎠EG

Thus the relative rates can be maximised by increasing temperature and increasing the water to solvent ratio. The temperature of the reactive zone of the catalytic distillation column should be kept as close to the Amberlyst 15 operating limit of 393 K as practicable. Beyond this operating limit, the ion exchange catalyst can be thermally deactivated. At atmospheric conditions this can be achieved through the use of a hot solvent and solvent feed temperature control. The powers of water and solvent terms, sets the requirement that the use of solvent in the catalytic extractive distillation be kept

Chapter 4: Kinetics of Hydration.... 155 to the minimum required for effective breakage of the azeotrope so water is maintained at the highest level possible to achieve required product purity.

1.0E-05 9.0E-06 341 K 8.0E-06 351 K )

-1 7.0E-06 361 K .s

-1 6.0E-06 5.0E-06

(mol.g 4.0E-06

MET 3.0E-06 r 2.0E-06 1.0E-06 0.0E+00 0 0.002 0.004 0.006

CIB (mol.L-1)

Figure 4.12 Formation rate of MET versus CIB for xEG = 0.81.

In summary, for the hydration of isobutylene in the presence of ethylene glycol basic treatment of the observed reaction rates suggested: • The hydration of isobutylene to TBA is pseudo first order with respect to isobutylene • The hydration of isobutylene in the presence of ethylene glycol is primarily affected through dilution of reactant water. Some non-linearity can be observed in the reaction rate with isobutylene as ethylene glycol concentration and temperature are increased • The use of ethylene glycol leads to a competing side reaction between isobutylene and ethylene glycol • That the rate of transport of isobutylene is sufficient for both the reaction of water and ethylene glycol with isobutylene

Chapter 4: Kinetics of Hydration.... 156

1.E-06 y = 1E-05x 1.E-06 2 R = 0.9337 ) -1

.s 8.E-07 -1 6.E-07 (mol.g 4.E-07 TBA r 2.E-07

0.E+00 0 0.02 0.04 0.06 0.08 0.1

CIB.CW (a)

1.E-05 9.E-06 8.E-06 )

-1 7.E-06 .s

-1 6.E-06 5.E-06

(mol.g 4.E-06

MET 3.E-06 r 2.E-06 1.E-06 0.E+00 0 0.02 0.04 0.06 0.08 0.1

CIB.CEG (b)

Figure 4.13 Rates of TBA and MET formation (a) water/isobutylene concentration product in the presence of ethylene glycol (b) ethylene glycol/isobutylene concentration product in the presence of water.

Chapter 4: Kinetics of Hydration.... 157 4.4 Heterogeneous Kinetic Models The rate of isobutylene hydration to TBA in the presence of ethylene glycol is modelled in terms of heterogeneous kinetics models. In this section relevant models are proposed, evaluated and scrutinised.

4.4-1 Proposed Models The following Langmuir-Hinshelwood mechanism for the reaction of water (W), isobutylene (B) to TBA (T) in the presence of ethylene glycol (E) was proposed:

WS+ ←⎯⎯⎯⎯→ WS.

B + SBS←⎯⎯⎯⎯→ .

WS..+ BS←⎯⎯⎯⎯→ TS .+ S

TS. ←⎯⎯⎯⎯→ T+ S

In addition to these steps, ethylene glycol is assumed to adsorb onto the same catalytic sites, thus reducing the number of free sites. Based upon this mechanism, the presence of ethylene glycol and assuming that the step of surface reaction was rate determining:

⎛⎞CT kCC⎜⎟WB− ⎝⎠K rT = 2 (4.68) ()1++++KCW W KC B B KC T T K EG C EG

Further, assumptions of insignificant product formation (CT ≈ 0) and KTBACTBA << 1 was made and the model of Equation (4.68) simplified to give:

kCBW C rT = 2 (4.69) ()1+++KCW W KC B B K EG C EG

The first variation of the model assumed that isobutylene is only weakly adsorbed or not adsorbed at all thus KBCB <<1.

Chapter 4: Kinetics of Hydration.... 158 The mechanism for ER type behaviour of isobutylene proposes that it reacts with adsorbed water directly from the fluid phase and is given as:

WS+ ←⎯⎯⎯⎯→ WS.

WS..+ B←⎯⎯⎯⎯→ TS

TS. ←⎯⎯⎯⎯→ T+ S

The Eley-Rideal model for isobutylene hydration and the assumption of competitive adsorption of ethylene glycol gives:

⎛⎞CT kCC⎜⎟BW− ⎝⎠K rT = (4.70) ()1+++KCW W KC T T K EG C EG

Accordingly, assuming little product formation (CT≈0) and KTBACTBA << 1, Equation (4.70) simplifies to:

kCBW C rT = (4.71) ()1++KCW W K EG C EG

Gonzalez and Fair (1997) in their study of the kinetics of isoamylene hydration treated swelling and reactant distribution by the incorporation of simple power laws to describe these, Equation (4.24). Therefore, while the rate of all physical transport steps is either excluded or accounted for in an effectiveness factor, information about the selectivity of these processes towards the components is retained. Considering specific distribution of reactants and allowing for association as well as adsorption, a power law model can be introduced incorporating the distribution coefficient and additional exponents. This treatment gives the general model given by:

lm kCIB C W rT = (4.72) mno ()1++KCWW KC EE

.

Chapter 4: Kinetics of Hydration.... 159

4.4-2 Model Evaluation The reaction rate data of isobutylene hydration in the presence and absence of ethylene glycol (Appendix A2, Table A2.03) were processed using non-linear regression using the software package POLYMATH 5. On account of the structure of the data, each temperature was treated separately and the results combined for further analysis. Additionally, this approach allowed for ease of evaluation in terms of thermodynamic considerations. Unified models were then proposed where possible and evaluated further in terms of statistical indicators and thermodynamic consistency.

The proposed models of Equation (4.66), (4.68) and (4.69) led to the development of seven primary heterogeneous kinetic models which are detailed in Table 4.10 along with the initial parameters evaluated. Four cases of the LHHW model were considered M1 to M4. The first, M1, assumed that all species adsorbed onto the same active sites, including isobutylene. This model failed to converge at all temperatures and was abandoned. It was modified assuming that isobutylene adsorbs weakly, if at all, and the

KIBCIB term was dropped from the denominator express. It was found that M2 did converge for all temperatures, however returned negative denominator coefficients, suggesting negative adsorption equilibrium constants. As these are considered rare it casts doubt on the validity of the M2. Subsequently, M3 and M4 assumed a power law distribution of water and ethylene glycol between bulk liquid and macro-pores. These models returned seemingly appropriate constants and coefficients. It was however found that the variation of exponents m and n in T were erratic for M3 and a unified model could not be developed for further analysis. M4 faired better and average values of m = 0.45 and n = 5.18 were found to give a stable solution. Of the LHHW set M2 and M4 were selected for further analysis and unified in temperature.

Next, three cases of the ER model (M5 to M7) were considered, the complete simplified Equation (4.68) and two power law modifications. The complete M5 was found to converge and provide reasonable reaction rate representation. Whilst the addition of the power law distribution terms for water and ethylene glycol gave marginally better representation. Again it was difficult to stabilise the model containing the “1” term in the denominator, ie. M6. However, M7 returned m = 0.6 and n = 5. The values of the

Chapter 4: Kinetics of Hydration.... 160 exponents appeared to be similar to those of the M4. Thus, M5 and M7 were selected for further analysis.

The temperature-unified versions of M2, M4, M5 and M7 were reported in Table 4.11. Three additional statistical indicators were employed in order to quantify the goodness of fit of the models:

• The Sum of Squares (SS) calculated as the sum of squares of the difference between the predicted and experimental values. The lower the SS the better the fit.

n 2 ˆ SS=−∑() rii r (4.73) i=1

• The Coefficient of Determination (COD) which equals the proportion of the variation in reaction rate that is explained by variables of the model. As COD approaches unity more of the variance of the data is accounted for.

nn 22ˆ ∑∑()rrii−− () rr ii − ii==11 COD = n (4.74) 2 ∑()rrii− i=1

• The F-Value, which takes into account the variance, accounted for by the model and

the number of parameters, np, is included in the model. A higher F-Value indicates a better, more reliable fit.

nn ⎡ 22⎤ ˆ ⎢∑∑()rrii−− () rr ii −⎥ ⎣ ii==11⎦

np F = n (4.75) 2 ∑()rrii− i=1

nn− p

Chapter 4: Kinetics of Hydration.... 161 The use of these parameters suggested that M5 and M7 gave the best fit returning the lowest SS, and COD very close to unity. The differences in SS and COD (4.58e-14 and 0.0177) were marginal. Parity plots for M5 and M7 are shown in Figure 4.14. However, both models had some tendency to overestimate the lower reaction rates observed. A comparison of F values showed that M5 (F = 4.9) was more efficient than M7 (F = 3.5), giving a better fit for the number of parameters employed. Thus, based upon statistical indicators, M5 was suggested to be the better model.

Qualitatively the parameters of M5 and M7 behave in a manner consistent with thermodynamic expectations. The rate constants increased with temperature. The equilibrium of adsorption constants were positive and decreased with temperature.

Based upon Equation (4.25) the ΔHad and ΔSad were determined and reported in Table 4.09. The test value according to rule 4: −Δ

Table 4.09 Enthalpy and Entropy of adsorption for Models M5 and M7

Model ΔHad,W ΔSad,W ΔHad,EG ΔSad,EG test W test EG (kJ.mol-1) (kJ.mol-1) (kJ.mol-1) (kJ.mol-1) M5 38.10 -0.050 38.76 -0.04 12.15 12.15 M7 23.40 -0.006 11.701 -0.057 12.17 12.18

The positive enthalpies of adsorption, indicating exothermic adsorption, were consistent with the decrease in magnitude of the equilibrium constants with increased temperature. The negative values of change in entropy were consistent with the principle that the molecules have less degrees of freedom in the adsorbed state and hence a lower level of disorder than in the fluid state. The Boudart, Mears and Vannice criterion returned a positive result well above the absolute of the change in entropy. Thus, both models may be considered thermodynamically consistent. However, it was also noted that the development of M5 has a stronger mechanistic basis.

Based upon both statistical and thermodynamic considerations M5 was selected as the best model. The ER mechanism has been employed previously for isobutylene reaction systems as evident in the literature review. M5 will be used in the evaluation of the reactive study.

Chapter 4: Kinetics of Hydration.... 162 Table 4.10 Heterogeneous Kinetic Models Development and initial Parameter Estimates

-1 -1 Model Source/Implication Form rTBA (mol.g .s ) T (K) Parameters Statistical Indicator R2 M1 LHHW complete - Failed to converge at all T - kCIB C W • adsorption IB 2 • adsorption W ()1+++KCIB IB KC W W K EG C EG • adsorption EG M2 modified LHHW 341 k = 4.73e-07 0.9539 kCIB C W • insignificant adsorption of 2 KW = -0.0120 IB ()1++KCW W K EG C EG KEG = -0.0272 0.9565 451 k = 2.44e-06 KW = -0.0082 KEG = -0.0149 0.9548

461 k = 8.29e-08 KW = -0.0194 KEG = -0.0458

M3 modified LHHW m 341 k = 9.763e-03 0.9546 kCIB C W • power law W and EG 2 KW = 2.496 mn m = 0.437 distributions between bulk ()1++KCW W K EG C EG liquid and macro-pores KEG = 1.887e-07 n = 6.942

451 k = 5.520e-03 0.9565 KW = 1.301 m = 0.427 KEG = 7.348e-06 n = 5.298

461 k = 4.045e-03 0.9804 KW = 0.731 m = 0.484 KEG = 6.693e-04 n = 3.480 Chapter 4: Kinetics of Hydration.... 163

M4 modified LHHW m 341 k = 0.183 0.9367 kCIB C W • power law W and EG 2 KW = 11.270 mn m = 0.45 distributions between bulk ()KCW W+ K EG C EG liquid and macro-pores KEG = 1.237e-06 • modified site balance n = 6.802

451 k = 0.1425 0.9564 KW = 7.539 m = 0.43 KEG = 4.768e-05 n = 5.22

461 k = 0.109 0.9540 KW = 4.661 m = 0.47 KEG = 3.271e-03 n = 3.52

M5 ER complete 341 k = 0.632 0.9780 kCIB C W • adsorption W KW = 2583.982 ()1++KC K C • adsorption EG W W EG EG KEG = 9741.154

451 k = 1.166 0.9197 KW = 2522.423 KEG = 10040

461 k = 1.688 0.8642 KW = 2224.7913 KEG = 10100

M6 modified ER m 341 k = 0.769 0.9546 kCIB C W • power law W and EG KW = 3247.420 mn distributions between bulk ()1++KCW W K EG C EG m = 0.781 liquid and macro-pores KEG = 1.428e-04 451 n = 7.776

k = 0.907 0.9563 Chapter 4: Kinetics of Hydration.... 164 KW = 1998.821 461 m = 0..649 KEG = 0.1245 n = 5.742

k = 0.522 0.9800 KW = 682.779 m = 0.34 KEG = 0.023 n = 4.87

M7 modified ER m 341 k = 5.391e-03 0.9546 kCIB C W • power law W and EG KW = 22.758 mn distributions between bulk ()KCW W+ K EG C EG m = 0.78 liquid and macro-pores KEG = 1.0e-06 • modified site balance n = 7.78

451 k = 5.61e-03 0.9565 KW = 12.363 m = 0.65 KEG = 7.699e-05 n = 5.74

461 k = 3.426e-03 0.9719 KW = 4.516 m = 0.60 KEG = 7.67e-03 n = 3.59

Chapter 4: Kinetics of Hydration.... 165 Table 4.11 Results of Heterogenous Kinetic Model Further Evaluation

-1 -1 Model Complete Form, rTBA (mol.g .s ) = Parameters Statistical Indicators 2 -1 -1 -1 M2 k0 = 6.83e24 (L .mol .g .s ) SS = 9.06e-13 ⎛⎞E -1 kCC0 exp⎜⎟IB W E = -203.79 (kJ.mol ) COD = 0.6461 RT ⎝⎠ a1 = -113.81 F = 3.3384 2 a2 = 0.3199 ⎛⎞⎛⎞a1 ⎛⎞a3 ⎜⎟1++⎜⎟aC24WEG ++⎜⎟ aC a2 = -405.53 ⎝⎠⎝⎠TT⎝⎠ a3 = 1.155 M4 k0 = 14.236 SS = 1.83e-13 ⎛⎞E m -1 kCC0 exp⎜⎟IB W E = -20.187 (kJ.mol ) COD = 0.9286 RT ⎝⎠ m = 0.45 F = 3.3661 2 n = 5.18 ⎛⎞⎛⎞a1 mn⎛⎞a3 ⎜⎟exp⎜⎟++aC24WEG exp⎜⎟ aC a1 = 2243.4 ⎝⎠⎝⎠TT⎝⎠ a2 = -5.561 a2 = 1382.5 a3 = -14.851 2 -1 -1 -1 M5 k0 = 345.65 (L .mol .g .s ) SS = 1.31e-13 ⎛⎞E -1 kCC0 exp⎜⎟IB W E = -19.5886 (kJ.mol ) COD = 0.9485 Adopted ⎝⎠RT a1 = 4582.6 F = 4.9008 a = -6.205 ⎛⎞⎛⎞a1 ⎛⎞a3 2 ⎜⎟1exp++++⎜⎟aC24WEG exp⎜⎟ aC a2 = 4662.2 ⎝⎠TT⎝⎠ ⎝⎠a3 = -5.0915 M7 k0 = 96446 SS = 8.52e-14 ⎛⎞E m -1 kCC0 exp⎜⎟IB W E = -34.548 (kJ.mol ) COD = 0.9662 RT ⎝⎠ m = 0.60 F = 3.5027 n = 5.00 ⎛⎞⎛⎞a1 mn⎛⎞a3 ⎜⎟exp⎜⎟++aC24WE exp⎜⎟ + aCGa1 = 2814.2 ⎝⎠TT⎝⎠ ⎝⎠a2 = -0.6562 a2 = 1407.3 a3 = -6.8064

Chapter 4: Kinetics of Hydration.... 166 1.2E-06

1.0E-06 ) -1 .s

-1 8.0E-07

6.0E-07

4.0E-07 model (mol.g TBA r 2.0E-07

0.0E+00 0 2E-07 4E-07 6E-07 8E-07 1E-06 1E-06 -1 -1 rTBA experimental (mol.g .s ) (a)

1.2E-06

) 1.0E-06 -1 .s -1 8.0E-07

6.0E-07

4.0E-07 model (mol.g TBA r 2.0E-07

0.0E+00 0.0E+ 2.0E- 4.0E- 6.0E- 8.0E- 1.0E- 1.2E- 00 07 07 07 07 06 06 -1 -1 (b) rTBA experimental (mol.g .s )

Figure 4.14 Heterogenous Kinetic Model Further Evaluation of (a) M5 and (b) M7 parity plots.

Chapter 4: Kinetics of Hydration.... 167 4.5 Concluding Remarks The main conclusion of this chapter regarding the hydration of isobutylene in the presence of the solvent ethylene glycol over Amberlyst 15 were: • The solubility of isobutylene was found to be approximately ten times higher in TBA than in water. Its solubility in ethylene glycol was found to be twice as high as for water. The solubilities within water and ethylene glycol were found to be approximately additive. The solubility within mixtures containing TBA was nonlinear and skewed towards that in pure TBA. The solubility of isobutylene within ternary mixtures of water, ethylene glycol and TBA was adequately represented by the O’Connell-Prausnitz model for a two way interaction mechanism:

lnHxHxHxH=+ ln ln +ln + IB... mix W IB W EG IB EG TBA TBA (4.43) ααW.. EGxx W EG++ W TBA xx W TBA α EG . TBA x EG x TBA where:

−03 αWEG. =×3.026 10T − 14.14 (4.44)

−03 αWTBA. =−7.281 × 10T − 14.88 (4.45)

−02 α EG. TBA =−2.069 × 10T + 3.53 (4.46)

• The activation energy of the forward reaction, pseudo first order in isobutylene in the absence of solvent was found to be 59.82 kJ.mol-1.

• The solvent ethylene glycol was not inert with MET as a by-product. The relative rate of TBA to MET formation could be described by the expression:

0.89 ⎛⎞−1127 ⎛⎞CW SRTBA=1.0exp⎜⎟⎜⎟ C W (4.67) ⎝⎠TC⎝⎠EG

• The following heterogeneous kinetic model expression based upon the LHHW formalism and an ER mechanism (Model M5, Table 4.15) provided an adequate representation of the kinetic data of the hydration of isobutylene:

Chapter 4: Kinetics of Hydration.... 168

⎛⎞−E kCC0 exp⎜⎟IB W ⎝⎠RT rTBA = (4.76) ⎛⎞⎛⎞a1 ⎛⎞a3 ⎜⎟1exp++++⎜⎟aC24WEG exp⎜⎟ aC ⎝⎠⎝⎠TT⎝⎠ where 2 -1 -1 -1 k0 = 345.65 (L .mol .g .s ) E = 19.5886 (kJ.mol-1) a1 = 4582.6 a2 = -6.205 a2 = 4662.2 a3 = -5.0915

The rotating basket reactor allowed for external mass transfer limitations to be eliminated. Thus if large particle sizes are used, dp 1mm, the reaction rate should only be limited by the diffusion within the particles. This improvement in external mass transfer will be returned to in Chapter 8, when a new catalytic distillation reactor is proposed, the Basket Impeller Column. This column uses the rotating basket here but with enhanced aspect ratio of impeller diameter to reactor diameter and higher catalyst loading. Unlike the current reactor which makes use of a sintered stainless steel plate which absolutely retains liquid upon gasification the column will use a sieve plate of much higher percentage open area and hole size. Thus deliberately allowing for weeping, creating a dualflow type flow scenario and achieving countercurrent flow.

Using the basket arrangement the kinetics of the hydration of isobutylene over Amberlyst 15 have been determined. These findings will be used to assess the extent of mass transfer in the pilot scale runs of the reactive runs within a countercurrent fixed bed reactor and a catalytic extractive distillation column. In addition to these reaction rates, a knowledge of mixing, retention time and holdup will assist in understanding the phenomenon taking place in these larger reactor systems of Chapter 7. Quite obviously the extent of mixing attained in the stirred high intensity reactor in this chapter is difficult to obtain within a column that relies on static geometry and flow due to the action of gravity. Thus the next chapter considers reactor fluidynamic characterisation of these static systems under different flow conditions.

Chapter 4: Kinetics of Hydration.... 169 4.6 Nomenclature

General a, b, c, d constants of correlations and models as defined in corresponding section or table A, B onstants used in Equation 4.53 and 4.54 ce cohesive energy density, J.mol.mL-1

Ci concentration of component i, (mol/L)

De effective diffusivity dp particle diameter (m) k reaction rate constant, units particular to model kj mass transfer coefficient, units particular to expression K reaction equilibrium constant

Ki adsorption equilibrium constant of component i ri reaction rate of component i (mol/g.s)

Rk resistance to mass transfer of step k, units particular to expression

sIB solubility od isobutylene (mol/L) smix solubility of gas in mixture sdev solubility deviation from ideality T temperature (K) vi volumetric fraction of component i xi mole fraction of component i Greek Φ Weisz Prater Criterion η effectiveness factor

ρi density of component i or species i Subscript / Acronym AS concentration of component A at surface S (mol/L) CED Cohesive Energy Density EG ethylene glycol

H2O water, generally represented as W

Chapter 4: Kinetics of Hydration.... 170 IB isobutylene MET ethylene glycol mono-tert-butyl ether TBA tert-butyl alcohol W water

Chapter 4: Kinetics of Hydration.... 171 4.7 Literature Cited Adesina, A. A. “Tools for mechanistic investigation: Application of the BMW-KD criterion to the methanation reaction.” J. Nigerian Soc. Chem. Eng. 7(1): 258- 263. (1988). Aiouache, F. and S. Goto “Sorption Effect on Kinetics of Etherification of tert-Amyl Alcohol and Ethanol.” Chemical Engineeering Science 58: 2065-2077. (2003). Boudart, M., D. E. Mears and M. A. Vannice Ind. Chim. Belge special issue: 281. (1967). Bowron, D. T., J. L. Finney and A. K. Soper “Structural investigation of solute-solute interaction in aqueous solutions of tertiary butanol.” J. Phys. Chem. B 102: 3551-3563. (1998). Boz, N., T. Dogu, K. Murtezaoglu and G. Dogu “Effect of hydrogen ion-exchange capacity on activity of resin catalysts in tert-amyl-ether synthesis.” Applied Catalysis A: General 268: 175-182. (2004). Carberry, J. J. "Chemical and Catalytic Reaction Engineering". NY, McGraw-Hill Book Company. (1976). Chaudhari, R. V. and P. A. Ramachandran “Three Phase Slurry Reactors.” AIChE Journal 26(2): 177-200. (1980). Delion, A., B. Torck and M. Hellin “Equilibrium Constant for the Liquid Phase Hydration of isobutylene over Ion-Exchange resins.” Ind. Eng. Chem. Process Des. Dev. 25: 889-893. (1986). Delion, A., B. Torck and M. Hellin “Hydration of Isopenetenes in an Acetone Environment over Ion-Exchange resin: Thermodynamic and Kinetic Analysis.” J. Catalysis 103: 177-187. (1987). Fisher, S. and R. Kunin “Routine Exchange Capacity Determination of Ion Exchange Resins.” Anal. Chem. 27(7): 1191-1194. (1955). Fogler, H. S. "Elements of Chemical reaction Engineering". New Jersey, Prentice Hall. (1999). Freda, M., G. Onori and A. Santucci “Infrared Study of the Hydrophobic Hydration and Hydrophobic Interactions in Aqueous Solutions of tert-Butyl Alcohol and Trimethylamine-n-oxide.” J. Phys. Chem. B 105: 12714-12718. (2001). Gates, B. C. and W. Rodriguez “General and Specific Acid Catalysis in Sulfonic Acid resin.” J. Catalysis 31: 27-31. (1973).

Chapter 4: Kinetics of Hydration.... 172 Gonzalez, C. J. and F. J. R. “Preparation of tert-Amyl Alcohol in a Reactive Distillation Column. 1. Reaction Kinetics, Chemical Equilibrium and Mass-Transfer Issues.” Ind. Eng. Chem. Res. 36: 3833-3844. (1997). Goto, S. and T. Saito “Liquid-solid mass transfer in basket type three=phase reactors.” J. Chem. Eng. Japan 17(3): 324-327. (1984). Gu, C. H., H. Li, R. B. Gandhi and K. Raghavan “Grouping solvents by statistical analysis of solvent property parameters: implications to polymorth screening.” Int. J. of Pharmaceuticals 283: 117-125. (2004). Gupta, V. P. and J. M. Douglas “Diffusion and Chemical Reaction in Isobutylene Hydration within Cation Exchange Resin.” AIChE Journal 13(5): 883-889. (1967). Hines, A. L. and R. N. Maddox "Mass Transfer, Fundamentals and Applications". New Jersey, P T R Rrentice Hall. (1985). Ihm, S. K., M. J. Chung and K. Y. Park “Activity Difference between the Internal and External Groups of Macroreticular Ion Exchange Resin Catalysts in Isobutylene Hydration.” Ind. Eng. Chem. Res.(27): 41-45. (1988). Jayadeokar, S. S. and M. M. Sharma “Ion Exchange Resin Catalysed Etherification of Ethylene and Propylene Glycols with Isobutylene.” Reactive Polymers 20: 57- 67. (1993). Kazanskii, V. S., S. G. Entelis and N. M. Chirkov Z. Fiz. Khim. 33(6): 1409-1413. (1959). Leung, P., C. Zorrilla, F. Recasens and J. M. Smith “Hydration of Isobutene in Liquid Full and Trickle Bed Reactor.” AIChE Journal 32(11): 1839-1847. (1986). Leung, P. C. “Solubilities and Enthalpies of Adsorption Of Isobutene into tert-Butyl Alcohol Water Mixtures.” J. Chem. Eng. Data 32: 169-171. (1987). Mears, D. E. Ind. eng. Chem. Process. Des. Dev. 10, 541 (1971) Myers, E. C. and K. K. Robinson ACS Sym. Ser. 65: 447. (1978). Njiribeako, A. I., P. L. Silveston and R. Hudgins Can. J. Chem. Eng. 56: 643. (1978). O'Connell, J. P. and J. M. Prausnitz “Thermodynamics of Gas Solubility in Mixed Solvents.” I & EC Fundamentals 3(4): 347-351. (1964). Parra, D., J. Tejero, F. Cunill, M. Iborra and J. F. Izquierdo “Kinetics study of MTBE liquid phase syhthesis using C4 olefinic cut.” Chemical Engineeering Science 49(24A): 4563-4578. (1995).

Chapter 4: Kinetics of Hydration.... 173 Pavko, A., D. M. Misic and J. Levec Chem. Eng. J. 21: 149. (1981). Prausnitz, J. M. "Molecular Thermodynamics of Fluid-Phase Equilibria". New York, Prentice-Hall. (1969). Rushton, J. H., E. W. Costich and H. J. Everett “Power characterisation of mixing impellers. Part I and II.” Chem. Eng. and Progress 46: 395-404, 467-476. (1950). Safinski, T. and A. A. Adesina “Development of a Novel Basket Impeller Dualflow Tray Catalytic Distillation Reactor.” Ind. Eng. Chem. Res. 44, 6212-6221 (2005). Velo, E., L. Puigjaner and F. Recasens “Inhibition by Product in the Liquid Phase Hydration of isobutene to tert-Butyl Alcohol: Kinetics and Equilibrium Studies.” Ind. Eng. Chem. Res. 27: 2224-2231. (1988). Velo, E., L. Puigjaner and F. Recasens “Intraparticle Mass Transfer in the Liquid Phase Hydration of Isobutene: Effects of Liquid Viscosity and Excess product.” Ind. Eng. Chem. Res. 29: 1485-1492. (1990). Yoshida, K., T. Yamaguchi, A. Kovalenko and F. Hirata “Structure of tert-butyle alcohol-water mixtures studied by RISM theory.” J. Phys. Chem. B 106: 5042- 5049. (2002). Zhang, C. M., A. A. Adesina and M. S. Wainwright “Isobutene hydration in a countercurrent flow fixed bed reactor.” Chem. Eng. and Processing 43: 533-539. (2003). Zhang, C. M., A. A. Adesina and M. S. Wainwright “Isobutene hydration over Amberlyst 15 in a slurry reactor.” Chem. Eng. and Processing 42: 985-991. (2003). Zhang, C. M., A. A. Adesina and M. S. Wainwright “Solubility Studies of Isobutylene in Tertiary Butyl Alcohol + Water Mixtures.” J. Chem. Eng. Data 47: 1476- 1480. (2002).

Chapter 4: Kinetics of Hydration.... 174 5. Bale Reactor Characterisation

The previous chapter provided information about the hydration of isobutylene in the presence of a solvent over Amberlyst 15. Small baskets were employed to hold the catalyst, while allowing for renewal of liquid immediately surrounding it. In this chapter the fluidynamic characteristics of a catalyst containment device, Bale packing, which allows for simultaneous reaction and separation, are explored. In the case of Bale packing the catalyst is retained in fibreglass bags rolled in a manner that provides channels for countercurrent flow of liquid and gas.

Although this packing is extensively used on an industrial scale there is little information regarding its mixing characteristics and open literature material regarding this topic has yet to be published. While random packing such as Raschig rings and structured packing have been significantly studied, their physical properties are very different and the correlations developed are not readily transferable.

An understanding of residence time, holdup and mixing allowed for better interpretation of reactive application of the packing under the CFBR and CED conditions of Chapter 7. Characterisation of holdup and pressure drop also provided opportunity for operational regimes to be identified which will assist in interpreting the effectiveness of Bale packing as an extractive distillation separation device.

This study was conducted in order to characterise an existing system of column diameter of 0.045 m. Experimental methods were developed to allow for the acquisition of RTD within this specific reactor. This chapter is divided into three parts, the first of which is concerned with basic capacity measurements performed manometrically and volumetrically. The second deals with gas phase mixing and residence times under ReG and ReL conditions as may be attributed to CFBR operation. The third part is concerned with liquid phase mixing in the preloading regime applicable to both CED and CFBR operation.

Chapter 5: Bale Reactor Characterisation 175 5.1 Background Firstly we provide a review of the relevant theoretical aspects related to packed column hydrodynamics. 5.1-1 Packed Column Capacity For operation of packed distillation columns, containing either random or structured packing, Kister (1992) describes four major hydrodynamic regimes:

™ The turndown maldistribution regime. Below a certain flowrate of liquid giving minimum wetting rate the falling liquid breaks up, some of the packing surface remains dry and efficiency is very poor. While gravity and viscous forces resist dewetting, the surface tension and vapour shear forces tend to dewet the falling film. The minimum wetting rate therefore rises with an increase in surface tension and liquid density, with a decrease in liquid viscosity. ™ The pre-loading regime. Most packed towers are designed to operate in this regime, which lies below 65 % flooding velocity (tray towers are operated above the loading point). Column efficiency is independent of flow rate and column pressure drops uniformly increases with vapour flowrate. The efficiency (E) of wire mesh packing and some corrugated sheet structured packing can decrease in this regime with both vapour and liquid flowrates. Raising vapour and liquid loads enhance mass transfer coefficients and interfacial area (E improves), but lower gas residence time (E decreases). Additionally, as liquid rate increases, more vapour is entrained down the column, lowering E. Structured packing permits far less lateral movement of fluids than random packing and more vapour is carried downward. ™ The loading regime. For all liquid flowrates a vapour velocity is reached when the gas velocity begins to interfere with free drainage of liquid. Liquid will start to accumulate or load the bed, hence the loading regime. Most columns can operate in this regime and achieve maximum efficiency. ™ The flooding regime. This regime is characterised by instability entrainment and poor efficiency and should be avoided. Further, increase in vapour velocity results in the packed column being inverted to a bubble column. The point of flooding indication, known as incipient flooding, can be very subjective and include symptoms such as appearance of liquid at the top of the bed, excessive entrainment,

Chapter 5: Bale Reactor Characterisation 176 a sharp rise in pressure drop, a sharp rise in holdup and a sharp drop in efficiency. However, fully developed flooding is described by: • the slope of the plot of velocity versus pressure drop goes to infinity • the gas velocity at which efficiency goes to zero

The regimes described can be most easily defined through basic capacity measures such as those of vapour pressure drop and liquid holdup. Various correlations and theoretical models have been developed to describe these within each regime of interest.

For a dry bed Kister (1992) showed that pressure drop could be calculated from a momentum balance and proposed the following form of correlation:

2 Δ=PCudry1ρ G S (5.01)

In order to allow for irrigation Kister modified the dry pressure drop correlation to give:

Cu2 L 2 Δ=PCwet 110 ρGS u (5.02)

This correlation applies to the gas continuous regime beyond which the pressure drop is no longer proportional to the square of vapour velocity. The transition from gas continuous to liquid continuous was said to occur at F = 0.4, where F is defined as:

L ρ F = G (5.03) G ρL

A more theoretical model, which has been applied for both random and structured packing, is that of the channel model. This model attributes pressure drop to resistance to flow in a multitude of parallel channels. Liquid flows down the walls of the channels consuming some of the available cross-sectional area thus increasing pressure drop. Bravo et al (1986) proposed the following for structured packing:

Chapter 5: Bale Reactor Characterisation 177 2 ⎛⎞214.3 ⎛⎞ρGeffu 1 Δ=P ⎜⎟0.395 + ⎜⎟ (5.04) Re ⎜⎟dg 0.5 5 ⎝⎠Geqc⎝⎠()1− CFr where u u = S (5.05) eff ()ε sinθ

dueq effρ G ReG = 124 (5.06) μG 12u2 Fr = L (5.07) dgeq

For Bale packing (dimensions d1 70 mm d2 200 mm and h 150 mm) arranged in an industrial manner with 7 bales per layer and 30o offset in rotation between layers, Xu et al (1997) obtained the following correlation for prediction of dry pressure drop:

ΔP dry =171.2U 1.790 (5.08) Z G and for wet bale packing below the flooding point:

⎛⎞ΔPwet 0.330 0.048 ln⎜⎟= 5.539UUGL (5.09) ⎝⎠Z

Equation (5.09) suggested that the influence of liquid velocity is much less than that of vapour velocity.

Stichlmair et al (1989) proposed a generalised fundamental model for pressure drop in packed towers. This model has been employed in the prediction of separation efficiency of Bale packing for the system cyclohexane/n-heptane 138 kPa by Subawalla, Gonzalez, Seibert and Fair (1997) and an adequate overall model fit was obtained. Dry pressure drop is expressed as a function of vapour velocity as:

Chapter 5: Bale Reactor Characterisation 178 ΔP 3 ⎡⎤(1−ε ) ρG 2 = f0 ⎢⎥4.65 UG (5.10) Zd4 ⎣⎦ε p where: AB f0 = ++3/2 C (5.11) ReGG Re

The constants of Equation (5.11) where evaluated for Bale packing (dimensions d 45 mm h 300 mm) packed one per layer with the use of three wall wipers per bale, Subawalla et al (1997) as: A = 753.896, B = -10.889 and C = 1.016. The irrigated wet pressure drop is given by:

(2/3+c) ⎡⎤11−−εεh / −4.65 ΔPwet ()()L ⎛⎞hL =−⎢⎥⎜⎟1 (5.12) Δ−P 1 εε dry ⎣⎦⎢⎥() ⎝⎠ where:

10.5⎡ AB⎤ c =−⎢ + 0.5 ⎥ (5.13) f0 ⎣ReGG Re ⎦ with A and B having the same values as for Equation (5.11).

Akbarnejad et al (2000) investigated the capacity of Bale packing (dimensions d 40 mm and h 200 mm) arranged in an industrial manner using 4 layers with 7 bales per layer and a 30o offset in rotation between layers. They attempted to use the generalised models of Eckert and that of Stichlmair et al (1989). There was however, no satisfactory agreement with either model. Thus they proposed a modified Eckert model. The dimensionless numbers X and Y for the original Eckert model are defined:

L ρ X = G (5.14) G ρLG− ρ and

Chapter 5: Bale Reactor Characterisation 179 0.1 2 G ⎛⎞μL YC= f ⎜⎟ (5.15) gρρGL()− ρ G⎝⎠ μ0

3 where Cf was defined for bale packing as a/ε . Isobars of pressure drop can be constructed using: ln()YABXCX=+ ln( ) + ln ( 2 ) (5.16)

The values of A, B and C corresponding to a particular ΔP/Z were tabulated, thus the model lends itself to the construction of a chart and needs to be reworked for use in general modelling and finer design calculations.

Liquid holdup is the liquid present in the void spaces of the packing. At flooding all the voids are filled with liquid or froth. Reasonable liquid holdup is necessary for good mass transfer and efficient operation and reaction in the case of CD. Overall liquid holdup is the sum of static and dynamic liquid holdup:

HHH0 = SD+ (5.17)

Static holdup, HS, is the liquid remaining on the packing after it has been fully wetted and drained for a long time. Its contribution to mass transfer is limited. Dynamic holdup, HD, is the liquid on the packing attributed to dynamic operation and is defined as the difference between total holdup and static holdup. Dynamic or operating holdup contributes to mass transfer as it provides residence time. Liquid holdup of packed columns can be measured by at least four different methods: ™ gravimetric, weighing the entire column/apparatus ™ volume, draining the packing after shutting off all inlet and outlet steams ™ tracer, RTD analysis for mean residence time from which holdup can be estimated via, Ramachandran (1983): tU H = m (5.18) 0 L

™ microwave, by an energy balance giving the amount of energy absorbed

Chapter 5: Bale Reactor Characterisation 180 Over Bale packing for the pre-loading regime Xu et al. (1997) obtained the following correlation:

0.0109 0.429 HuuDL, = 0.0336 GL (5.19)

indicating the minor influence of uG. Xu et al labelled the pre-loading, loading and flooding regimes as the film, bubble and emulsification regimes respectively. They also observed some degree of hysteresis in their measurements especially around the transition between pre-loading and loading regimes.

For liquid holdup of bale packing Subawalla et al (1997) used the generalised model of Stichlmair et al (1989). In this model total liquid holdup is seen as the sum of that due to wall or film flow and that due to packed flow in the pre-load regime given as:

HH= wall+ H pack (5.20)

where Hwall results from a momentum balance of liquid flowing down an inclined plane:

1/3 ⎛⎞12 2 Hauwall = ⎜⎟wallν L L (5.21) ⎝⎠g

and Hpack results from adaptation of Equation (5.21) in order to account for porosity effects:

1/3 ⎛⎞12 2 uL Hapack= ⎜⎟ packν L 4.65 (5.22) ⎝⎠g ε for use with Bale packing the wall contribution stems from both the column wall and the external cloth surface.

Akbarnejad et al (2000) proposed the following correlation for HD for bale packing operated in the pre-loading regime:

0.39 HLDL, = 0.042 (5.23)

Chapter 5: Bale Reactor Characterisation 181 where L is the liquid mass flux (kg.m-2.s-1).

5.1-2 Reactor Fluid Dynamics and Residence Time Distribution The mixing properties of bale packing have yet to be published in open literature. The mixing properties of other forms of catalytic packing have begun to be investigated with studies conducted over Sulzer Chemtech KATAPAK-S, Moritz and Hasse (1999), Ellenberger and Krishna (1999). Their findings indicated high degrees of dispersion, due to the multiple levels of porosity of the open form catalytic packing. They used the axial dispersion model for the characterisation of the packing backmixing. This section describes the RTD analysis methods and basic one-dimensional models for RTD interpretation and reactor characterisation.

RTD data resulting from an impulse concentration disturbance are typically normalised giving the exit age distribution: Ct( ) Et()= ∞ (5.24) ∫ Ctdt() 0

The first moment of which is mean residence time, tm:

∞ ttEtdt= (5.25) m ∫ () 0

The second moment gives variance, σ2:

∞ σ 2 =−tt2 Etdt (5.26) ∫()()m 0

The third and fourth moments are skewness and peakedness respectively, however error accumulation makes their estimates relatively poor. The method of moments suffers

Chapter 5: Bale Reactor Characterisation 182 from tailing effects in the same manner as averages suffer from outliers and thus it is often used as a first estimate.

The Dispersion and Tanks in Series models are well defined and explained in the reaction engineering texts of Fogler (1999) and Levenspiel (1972). Briefly discussed here are practical issues relating to the application, use and comparison of the dispersion model. The Stagewise Backmixing model, a basic two parameter model, with practical attributes is also introduced and discussed.

The dispersion model stems from a flow analogy to Ficks law, which results in the differential equation given by:

∂∂2CC∂ (UC) D −= (5.27) a ∂∂∂zzt2

where Da is the eddy (or axial or longitudinal) diffusion (dispersion) coefficient. The Reactor Peclet number, is defined as:

UL Per = (5.28) Da where U is the average or superficial velocity (U = Q/A) and L is the reactor/bed length. At this point it was considered necessary to introduce other definitions of Peclet number in use to allow for comparison. The Peclet number, also referred to as the fluid Peclet number, Pef, is more rigorously defined as:

ud p Pe f = (5.29) Da where u is the interstitial packed bed velocity defined:

U u = (5.30) ε packing

Chapter 5: Bale Reactor Characterisation 183

Bodenstein number, Bo, is also used and is defined as:

uL Bo = (5.31) Da hence sharing elements of the two previous definitions. In a similar manner Re can be defined as superficial Re, interstitial Reu and in between ReB (see Nomenclature). An understanding of which is being used is important for purposes of comparison and final application.

The boundary conditions used in the solution of Equation (5.27) are important in packing and reactor mixing evaluation. The closed-closed BC, or Danckwerts BC, assume that there is no dispersion before or after the reactor/defined length:

−+ at zz== 0 Daa 0 and ==L D 0

The closed-closed BC are typically used with impulse/step concentration disturbance at entrance and response analysis at reactor/packed length exit. Application of these BC has been shown to give:

2 σ 22 −Per 22=−()1 −e (5.32) tPePemrr

The open-open BC assume that dispersion exists before and after the defined length:

−+ at zz= 0 Daa>=> 0 and L D 0

The open-open BC are typically used with the two-point analysis experimental procedure popularised for packing evaluation/characterisation, Marcias-Salinas and Fair (2000). The claimed advantage is that end effects can be eliminated and a true measure of packing dispersion is obtained. Application of these open-open BC has been shown to give:

Chapter 5: Bale Reactor Characterisation 184 σ 2 28 22=+ (5.33) tPePemrr

A practical analytical solution to Equation (5.27) for C/Ctotal = f (Per, ( t/tm )) is available in dimensionless form for open-open BC and thus, two-point experimental investigation RTD data can be regressed directly against it yielding Per and tm. The solution closely resembles the normal distribution found in statistics.

Equations (5.32 & 5.33) can be rearranged and solved for Per by iteration using tm and σ2 obtained from method of moments. The accuracy of the second moment suffers from the effects of tailing and choice of data cut off point. More advanced methods include transforming the data with say the Laplace transform and carrying out the analysis in the frequency domain, [Mecklenburgh and Hartland (1975)]. These methods allow for the suppression of tailing effects. Mecklenburgh and Hartland (1975) showed that for unsteady injection at the inlet, solution of the dispersion model gives the following transfer function:

x = xgTin ( ) (5.34) where T is the Laplace Transform parameter and

⎡⎤⎛⎞Perr(11−−− z) ⎛Pe() z ⎞ ⎢⎥()1+−aa exp⎜⎟() 1− exp ⎜ ⎟ ⎛⎞Pe z 22 gT = 2exp r ⎢⎥⎝⎠ ⎝ ⎠ (5.35) () ⎜⎟⎢⎥ ⎝⎠2 22⎛⎞Perr ⎛−Pe ⎞ ⎢⎥()1+−−aa exp⎜⎟() 1 exp ⎜ ⎟ 22 ⎣⎦⎢⎥⎝⎠ ⎝ ⎠ and

4T a =+1 (5.36) Per The function g(T) can be simplified given particular experimental conditions and extent of backmixing. The parameter T can be set to suppress the effect of tailing while retaining the characteristic shape of the exit age distribution in the transformed domain.

Chapter 5: Bale Reactor Characterisation 185 5.2 Bale and Raschig Ring Packing In this section are reported the fundamental properties and preparation procedure of the laboratory scale Bale packing used throughout this study. Also described are the properties of 8 mm Pyrex Raschig rings. This small-scale first generation random packing was found to have approximately the same porosity as the Bale packing prepared and was used as benchmark packing throughout the fluid dynamics study.

5.2-1 Bale Packing Preparation Procedure The Bale Packing consists of a fibre glass bag containing many pockets of catalyst, which is wrapped with wire mesh to form a bale, Figure 5.01. Smith (1980) and Subawalla (1997) have described its fabrication. To construct the fibreglass bags, a locally available fibreglass cloth was sought. Without the availablity of a specialist industrial sewing machine the cloth had to be able to be processed on a home sewing machine. It was found that glass mat and most simple weaves such as plain, twill, four- shaft satin and crows foot were inappropriate. These weaves suffered from bunching of fibres upon sewing with a variety of stiches, needle and thread sizes. This bunching produced holes in the final bag, which allowed for the escape of particles in the 400 to 600 μm range. It was found that eight-shaft satin worked well with polyester thread. This cloth had a density of 127 kg/m3 and was kindly donated by Colan Products Pty Limited (Product No AF510HT127).

The fibreglass bags containing Amberlyst 15 or glass particles were fabricated by:

1. Cutting a rectangle 240 × 300 mm of cloth. 2. Folding over the rectangle to gain one 120 × 300 mm. 3. Sewing one (120 mm) end from fold to opening. Sewing runs at 3mm, and at 13 mm from this initial run and repeating this pattern. In this manner pockets of 13 mm width spaced 3 mm apart were formed. The quilt was formed of 16 pockets and the excess in terms of length cut off. 4. The bags were then filled with Amberlyst 15 using a small glass funnel and a glass rod. The pocket was first opened by inserting the glass rod and then the funnel was installed and Amberlyst poured in up to a pre-drawn 100mm guide line. The pockets

Chapter 5: Bale Reactor Characterisation 186 containing catalyst were massaged so as to remove voids and achieve good packing and complete pocket roundness. 5. Once each pocket had been filled and all were readjusted for uniformity, the top of the bag was sewn shut creating a 100 × 260 mm catalytic quilt.

Stainless steel mesh in 304-grade stainless steel, with 0.5 mm aperture and 0.315 mm wire diameter was sourced from Melwire Pty Limited (PN23177620). Strips were cut of 100 × 350 mm with a pair of tin snips.

Figure 5.01 Laboratory Scale Bale packing prepared.

The mesh and the bag were combined and rolled with the bag on the inside. An effort was made to obtain tight uniformly rolled bales. A piece of brass pipe with internal diameter of 40 mm was used as a guide. Excess wire mesh was trimmed off and the bales tied with a strand of the wire of the mesh by twisting the wire with a pair of pliers. Each bale was then equipped with three wall wipers the first located at the top of the bale (indicated by the sewn end of the bag). The next two were evenly spaced below. A wall wiper consisted of wrapping a thin layer of teflon plumbing tape around the bale,

Chapter 5: Bale Reactor Characterisation 187 then using a 5 mm strip of teflon plastic which was wrapped around the bale and trimmed to size and finally sealing this with a layer of teflon tape. The thickness of the tape was adjusted to give a snug fit in the reactor of internal diameter 45 mm.

Five Bales were prepared containing Pyrex spheres of diameter 853 to 1003 μm, representing the average size of the swollen resin. A further ten Bales were prepared containing Amberlyst 15 in the size range 500 to 699 μm.

5.2-2 Packing Characterisation The fundamental characteristics determined included porosity, catalytic loading, and static holdup. The characteristics were determined using volume displacement and gravimetric methods. Porosity and static holdup of water upon the packing at 298.15 K are reported in Table 5.01 While of similar porosities Bale packing retained substantially more static liquid, as may be expected of its bag like structure and the presence of small tighly packed particles. The bale packing was found to contain an average mass of Amberlyst of 37 g/bale and thus five bales gave a total reactor loading of 185 g or 233 g/L.

Table 5.01 Porosity and Static Holdup of Bale and Raschig Packing.

Packing Stage Particle HL, Static HL, Static Porosity vol. frac. vol. frac. vol. frac. vol. frac. Instant Rested Bale Packing Glass 850 to 1003 μm 0.7171 0.1886 0.2203 0.2047 Amberlyst 500 to 700 μm 0.7485 0.1572 0.2940 0.2578 Empty, Bag and Mesh 0.9057 na 0.0965 0.0884 Raschig Packing 8 mm 0.7257 na na 0.0536

Chapter 5: Bale Reactor Characterisation 188 5.3 Capacity of Bale Packing The experimental evaluation of pressure drop and liquid holdup over Bale packing was conducted within a pilot scale column, of diameter 45 mm, using water and air at ambient temperature and pressure. The range of flowrates used was chosen to represent those typically used for distillation. The results were correlated against relevant variables and compared against those of similar studies. The study was short and simple and its range and depth were extended in later sections using dynamic methods of collection and interpretation of residence time distributions. 5.3-1 Experimental The reactor or reactive section was a jacketed glass column fitted with ground glass quick-fit connections at either end of internal diameter 45 mm and external jacket diameter 67 mm, Figure 5.02. It was fitted at the top with a liquid distributor/gas collector and at the bottom with a gas distributor/liquid collector the details of which are given in Figure 5.03. The column was mounted vertically, levelled and secured within a heavy steel experimental cupboard.

Pressure drop was measured directly by the use of pressure gauges of range 0 to 250 kPa with limit of reading ± 1.0 kPa. Liquid phase holdup was measured using the volumetric draining technique, Akbarnejad et al (2000).

The rig used for these runs was very similar to that of Section 5.5.1.1. depicted in Figure 5.17. Water at ambient temperature was recirculated by means of pump from the collector down through the packing. A flowmeter was used to check for constant rate of recirculation (range: 0.03 to 0.2 L.min-1). The gas used was building utility air supplied at 600 kPa reduced to 250 by regulator and controlled by a large needle valve equipped with vernier scale (range: 15 to 95 L.min-1). The gas was saturated with water before being fed to the CFBR by passing it through a 2 L stainless steel vessel containing Raschig Ring packing and water. This ensured minimal loss of liquid due to stripping. The liquid collector equipped with graduated level sight was calibrated (range: 0 to 0.55 L), thus allowing for the determination of the distribution of liquid between collector and packing.

Chapter 5: Bale Reactor Characterisation 189

A typical run consisted of:

• taking the initial liquid collector level reading of the completely drained packing • setting the desired liquid and gas flow rates gradually • monitoring the flowrates and allowing for steady state to be reached • recording the pressure drop • taking down the new collector liquid level every 5 minutes for at least 30 minutes

Chapter 5: Bale Reactor Characterisation 190

G

Condenser

L L Distributor

id 45 mm

od 67 mm

G distributor Sampling Window

G

L

Figure 5.02 Detail of Reactive Zone Jacketed Reactor and distributors

Chapter 5: Bale Reactor Characterisation 191 900

tappered end QF 19/26 50/42

d10 550

d2"

d1/4"

QF 50/42 460 d1/4" mesh d20 nozzle 400

d1/2" 50

base support

(a) (b)

Figure 5.03 Detail of (a) top Liquid Distributor and Gas Collector fabricated in Pyrex and; (b) bottom Gas Distributor and Liquid Collector fabricated in brass and copper

Chapter 5: Bale Reactor Characterisation 192 5.3-2 Results and Discussion

5.3-2-1 Pressure Drop As expected pressure drop across the bed of Bale packing increased monotonically with

ReG and a small increase was observed upon wetting and increased ReL, Figure 5.04. Pressure drop measurements suffered from liquid entrainment and hence the noisiness of the curves obtained.

60 dry 50 ReL 15 ReL 43

) ReL 72 -1 40 ReL 100 30

P/L (kPa.m 20 Δ

10

0 0 1000 2000 3000

ReG

Figure 5.04 Pressure Drop of Dry and Wetted Bale Packing

The observed pressure drops of dry and wetted Bale were well described by the form of correlation proposed by Xu et al (1997). The dry pressure drop was adequately described by ΔP = 53.8836U 1.707 (5.37) L G with an R-square value of 0.9881, while the wetted pressure drop given by:

Chapter 5: Bale Reactor Characterisation 193 ΔP ln= 4.957UU0.468 0.023 (5.38) L GL has R squared = 0.9616.

Both Equations (5.38) and (5.39) gave qualitative agreement with that of Equation (5.08) and (5.09) proposed by Xu et al (1997). The wetted pressure drops returned a deviation of less than 3 kPa for the comparison to their published correlation.

5.3-2-2 Liquid Holdup Liquid holdup presented in Figure 5.05 varied little with gas flow rate in the pre-loading regime denoted by the flat portions of the family of curves. The steep increase of holdup beyond a certain point for each curve represented the loading regime. For bale packing there are few observable phenomena suggesting loading. The wall wipers tend to play a greater role and liquid temporarily accumulates at these before being reintroduced into the bale. The loading regime was narrow and rapidly followed by the onset of flooding. Flooding was observed to start at the base of the column and progressed slowly up its length marked by a transition to a continuous liquid phase and bubble column type appearance, which eventuated in considerable froth collection at the top of the column.

Chapter 5: Bale Reactor Characterisation 194 0.4

0.35

0.3

0.25

0,L 0.2 H 0.15 ReL 15 ReL 43 0.1 ReL 72 0.05 ReL 100 0 0 1000 2000 3000

ReG

Figure 5.05 Liquid Phase Holdup of Bale Packing

In the pre-loading regime (cf. Loading Equation (5.42)) liquid holdup did, however, increased with liquid velocity. Total holdup was correlated using:

0.27 HU0,LL=+0.22 0.04 (5.39) or in terms of ReL:

−03 0.27 H0,LL=+×0.22 2.1 10 Re (5.40)

The intercept corresponds to the static holdup, HS and was estimated through extrapolation of the data. This value compares reasonably well to that of 0.22 determined gravimetrically for bale characterisation, Section 5.2. The difference may be accounted for by the presence of wall wipers, which marginally increased static holdup. The power law term defines dynamic holdup which varies as a function of liquid velocity alone.

Chapter 5: Bale Reactor Characterisation 195 It can be seen that the regimes are clearly demarcated by changes in holdup behaviour and the onset of dependence upon UG, while there is little difference to be observed between these two regimes for pressure drop measurements. The loading and flooding envelopes defined by holdup measurements were presented in Figure 5.06.

3000 Loading Point 2500 Flooding Point 2000 G 1500 Re 1000 500 0 0 20 40 60 80 100 120

ReL

Figure 5.06 Loading and Flooding Points for Countercurrent operation of Bale Packing

The curves were most adequately described by:

−0.78 ReG, Loading = 21580ReL (5.41) and

−0.35 ReG, Flooding = 7357 ReL (5.42)

for the range 500 ≤ ReG ≤ 3000 and 10 ≤ ReL ≤ 100. These curves along with further operational regime data gained in the gas and liquid phase fluid dynamics studies allowed for better understanding of the CFBR and CED systems, in terms of flow regime.

The holdup associated with loading and flooding, Table 5.02, can be compared to the porosity of the packing found to be ε = 0.7171, Table 5.01, basic bale characterisation. It can be seen that in the preloading regime liquid occupies about 25 % of the free

Chapter 5: Bale Reactor Characterisation 196 packing space. As flooding is approached and just before the liquid becomes continuous it can be seen that it occupies close to 50 % of the free space.

Table 5.02 Fraction of Bale packing porosity occupied by liquid in the pre-loading regime and close to the flooding regime.

ReL H0,L H0,L/ε H0,L close H0,L/ε pre-loading to flooding 15 0.22 0.31 0.31 0.44 43 0.23 0.32 0.38 0.53 72 0.25 0.34 0.36 0.51 100 0.26 0.36 0.32 0.45

Considerable gas flow rates are required to achieve the loading regime over this range of liquid flow rates. In terms of steam the required flowrates were calculated and reported in Table 5.03, along with the reboiler duty required assuming no losses.

Table 5.03 The corresponding flowrates of water and steam (saturated 373 K) to

the loading ReG and ReL obtained and the ideal reboiler duty required to achieve the steam flowrates

ReL QL ReG QG q ideal (L.min-1) (L.min-1) (J/s) 15 0.0276 2338 100 2242 43 0.0789 1370 59 1314 72 0.1302 815 35 782 100 0.1816 502 21 481

Steam was considered as it is to be used in the CED runs and is the primary reactant. With a typical laboratory mantle of maximum power of 1200 Watts compensated with a top mantle of 1200 Watts and assuming 60 to 80 % efficiency, duties of 1440 to 1920 Watts may be attainable. Thus, operation was considered possible within the preloading regime, the transition between regimes and some runs could be achievable within the loading regime. In the next two sections liquid and gas phase mixing are considered primarily within the pre-loading regime.

Chapter 5: Bale Reactor Characterisation 197 5.4 Gas Phase Mixing, CFBR Flow Conditions: Pre- Loading Regime It was considered useful to determine gas phase residence time and mixing over Bale packing and Raschig ring packing within the range of flow rates of gas and liquid typically used for Countercurrent Fixed Bed Reactor (CFBR) operation because: ™ Under low gas flow rates for a relatively short bed the assumption of plug flow may not be adequate. This assumption is commonly made for countercurrent processes such as CD where vapour flow rates are relatively high and overall column lengths are long ™ While a handful of studies have been conducted regarding mass transfer phenomena over Bale packing there is no data concerning the underlying mixing characteristics or residence time ™ An understanding of gas phase mixing will aid reactor modelling and results interpretation

In addition to Bale packing, Raschig packing was evaluated because: ™ Some information on residence time distribution and holdup is available for Raschig ring packing, and thus, the current study and methods being developed could be bench marked against existing work ™ Raschig rings themselves can be manufactured of catalytic material and used within a CFBR or CD column, Kunz and Hoffmann (1995) ™ 8 mm Raschig rings were chosen for use in the non-reactive sections of the CED column.

The aims of this section were to: ™ Develop an experimental technique capable of acquiring gas phase RTD of an existing reactor. ™ Analyse the RTD acquired in a meaningful manner which did not corrupt the inherent information ™ Apply suitable models to characterise the observed fluidynamic effects and to correlate the dependent variables obtained with suitable variables allowing for reactor characterisation

Chapter 5: Bale Reactor Characterisation 198 5.4-1 Experimental, Gas Phase The RTD of the gas phase was obtained and analysed assuming that end effects were negligible compared to the effect of the packing. The ratio of reactor diameter d to bed length L was 1:13 and the reactor ends were of relatively simple and non-obstructive design. The ratio of Raschig Ring packing to diameter was 1:6, which is just below the minimum recommended 1:8 for the avoidance of wall effects. However, practical considerations such as better combination with Bale packing and the possibility of fabrication of catalytic rings of the same size promoted their use. Further the size required to completely avoid wall effects for a 45 mm diameter column of packing of diameter of 1 to 2 mm was obviously impractical. A diameter of 45 to 50 mm was required to allow for sufficient vapour velocities to be gained using small-scale 240 V reboilers and prevent excessive use of resources. The Bale packing was fitted with three wall wipers per bale, thus avoiding wall effects.

Pulse injection at inlet and sampling at exit were adopted as the experimental approach. Experience with pressure drop measurement showed that sampling vapour without significant liquid entrainment is difficult to achieve. Given the desired use of GC-TCD for response acquisition the avoidance of liquid and blockage was considered crucial and thus external sampling was adopted.

In order to observe the rapid transient behaviour of the gas phase dynamics automation and computer aided data acquisition was employed. Gas phase RTD measurements were made following a 20 mL pulse injection of helium tracer into flowing nitrogen with the aid of an actuated 6-way Valco valve. The experimental configuration is shown in Figure 5.08. The exit stream from the CFBR was split into two with a measured controlled reproducible flow diverted to an isothermal, thermal conductivity detector (Shimadzu 8A GC). The continuous output voltage signal of the GC was acquired using a National Instruments DAQ Card 6035E running with LabVIEW TM software. Actuation of the 6-way valve was computer-controlled and was used to trigger the data acquisition at 12 to 25 samples per second. Once gas and liquid flows were stabilised the TCD was dynamically zeroed.

Chapter 5: Bale Reactor Characterisation 199

5.4-1-1 Automation and DAQ with LabVIEW The front panel of the LabVIEW program, Blue.vi, employed for automation and data acquisition is shown in Figure 5.07. The extension vi stands for virtual instrument. The program Blue.vi was created through modification of the LabVIEW example Cont(inuous) Acq(uire) to Spreadsheet File.vi. The program could continuously display the voltage signal from the GC allowing for dynamic zeroing and for the experiment to be followed in real time. Blue.vi used a circular buffer technique of data acquisition whereby data is continuously acquired into a circular acquisition buffer at the same time that the program reads the acquired data and writes it to file. The rate of data acquisition could be set in the range 1 to 250 Hz. The program enabled a relay to be activated which switched an auto injector and produced a trigger which signalled data acquisition to be commenced. The acquired data was stored in a text file. When the program was placed in stand by, it requested the name of the intended text file. When the Boolean Start button was pressed by using the mouse, the injection was made and the data acquisition commenced till the Stop button was pressed. Acquisition was stable and free of hardware based disruption.

Figure 5.07 Front Panel of the Program Blue.vi used for Automation and Data Acquisition for Gas Phase RTD Experiments

Chapter 5: Bale Reactor Characterisation 200

vent [05] TCD vent [06]

[01]

[02] [07] [04]

[03] [08] [09] R R Water Tank

[10] He N2 Drain

Figure 5.08 Rig for Gas Side Fluid Dynamics Study

Key to Figure 5.08 CFBR Gas Side Fluid Dynamics [01] Liquid Flow Meter and Controller, BASIS Micro-Motion M/N F025S1319SM and Eurotherm 2216 PID controller. [02] Air actuated control valve and current to pressure transducer, Fisher Controls, M/N 912346. [03] Centrifugal pump, Grundfos, CH12-40-A-W-6-BQQV, 640 kW, h 37m. [04] Trickle Bed Reactor jacketed column. [05] Sample stream split rotameter and flow adjustment.

Chapter 5: Bale Reactor Characterisation 201 [06] TCD, Thermal Conductivity Detector, Shimadzu GC-8A operated withou column. [07] Computer with data acquisition card, National Instruments DAQ Card 6035E, and digital I/O card running relays. Controlled with LabVIEW program. [08] Injection valve, Valco, and automatic pneumatic drive. [09] Nitrogen gas supply and rotameter and Helium tracer supply. [10] Injection loop.

5.4-1-2 CFBR Gas Side Fluid Dynamics Operating Procedure 1. Open utility air to charge various actuators, flow control and autoinjector. 2. Set gas flow rate as required using the gas side rotameter [9]. 3. Adjust the sample stream split [5], providing a measured flow to one side of the TCD. 4. Turn on the GC 8A making sure that the primary carrier gas nitrogen is on and the other side of the TCD registers. Allow 1 hour for stabilisation and set the TCD current. 5. Set the liquid flowrate using the controller [01], allow time for steady state to be achieved. Note that the liquid level of the liquid collector below the CFBR must be maintained to prevent gas short cutting by the drain. Note that the packing needs to

be wet for ReL = 0 runs. 6. Fill the tracer injection loop [10] with Helium using the on off gas valve. Allow the Helium to keep flowing through the loop. 7. Turn on the computer and run the LabVIEW program, Blue.vi. Zero the base line of the GC by monitoring the signal and adjusting the detector output at the GC. 8. Check the system is ready for injection, that all flowrates are correct and that the liquid level is stable. 9. Inject the tracer by pressing start. Monitor the experiment press stop once sufficient time has elapsed and the signal has returned to the base line.

Chapter 5: Bale Reactor Characterisation 202 5.4-1-3 Experimental Design This section describes the choice and range of experimental variables used in the investigation of gas phase holdup and mixing properties over packings of interest to the CFBR and CED based hydration of isobutylene in the presence of the solvent ethylene glycol. Two packing types were investigated: 1. Bale Packing containing glass spheres of 850 μm with properties of d = 0.043 m, h = 0.1 m, porosity ε = 0.7171. 2 -3 2. Raschig ring packing of diameter dp 0.008 m, area ap 2027 m .m , porosity ε = 0.7257 fabricated of Pyrex glass and packed randomly.

The effect of gas phase Reynolds number upon gas phase holdup and mixing was investigated over such a range as may be used in a CFBR. Under CFBR conditions isobutylene in the gas phase may be fed in the presence of an inert diluent of low solubility such as n-butene in the final application and nitrogen for the purpose of experimental investigation. This range is typically different from that for CD where high vapour flow rates are involved due to the presence of boiling and where typically the assumption of plug flow conditions is made Taylor and Krishna (2000).

Zhang, Adesina and Wainwright (UNSW (2003)) for the same CFBR using however -1 -1 basket type packing employed a gas velocity of UG 0.189 cm.s (QG 0.18 L.min ) and -1 -1 liquid velocities of UL 0.44 to 2.52 cm.s ( QL 0.42 to 2.4 L.min ). They assumed plug flow in gas and found mass transfer to be dominated by liquid phase mixing. Alicilar et al (2002) investigated catalytic wet air oxidation of cyanide in a laboratory CFBR of d

56 mm, L 640 mm packed with 3 to 5 mm pellets. They employed gas velocities of UG -1 -1 -1 0.31 to 0.94 cm.s (QG 0.3 to 0.9 L.min ) and liquid velocities of UL 1.57 to 4.72 cm.s -1 (QL 1.5 to 4.5 L.min ), operating in the full range of the pre-loading regime.

CFBR experiments were planned in the preloading regime with low to moderate gas flow rates. Such flowrates give ample opportunity and sufficient mean residence time for good use of sparingly soluble gas phase reactant, such as isobutylene, over short catalytic beds. With literature suggesting both ReG and ReL influence gas phase holdup and mixing and a well-established science of fluid dynamics a complete matrix type experimental design was adopted. Gas Reynolds number was varied as follows:

Chapter 5: Bale Reactor Characterisation 203 ReG : 8, 20, 56, 92 and 129 -1 UG (cm.s ): 0.3, 0.7, 2.0, 3.3 and 4.5

While comparison to mixing data over Raschig ring packing as used for absorption studies would have been better made at higher gas flow rates, it was found that the helium tracer injection mechanism could not be operated beyond ReG 129.

The effect of liquid phase Reynolds number upon gas phase holdup and mixing was explored to the point of flooding. For low gas flowrates the loading regime was very narrow and could not be distinguished from the flooding regime. Preliminary liquid phase RTD runs showed invariance of H0,L in ReG over the range of the study. Further preliminary pressure drop measurements over Bale packing at increments of ReG = 30 and ReL = 123, helped define the region of stable operation. The pressure drop measurements of Table 5.04 were made with manometer using water and connected differentially across the bed. The transition to loading and flooding over Bale packing was defined by:

−1.109 ReGB..= 580633ReLB (5.43) Thus, the data was collected entirely in the pre-loading regime with Reynolds number varied as follows in the complete matrix design:

ReL : 0, 245, 490, 735, 981, 1226 and 1471 -1 UL (cm.s ): 0, 0.5, 1.0, 1.6, 2.1, 2.6 and 3.1

Table 5.04 Bale packing Pressure Drop, ΔP/L (kPa.m-1)

ReG: 188 218 248 277 307 337 366 396 426 456 -1 ReL ΔP/L (kPa.m ) 0 0.39 0.62 2.70 2.87 3.04 4.56 6.42 8.78 11.48 14.35 245 0.39 0.57 1.54 2.70 8.44 10.81 20.26 13.51 30.39 38.83 368 0.44 0.69 2.19 5.57 8.78 14.01 25.33 47.27 57.40 60.78 490 0.47 0.78 2.53 3.88 6.25 10.81 13.84 18.57 27.01 45.59 613 0.47 0.83 2.36 6.25 11.82 20.26 25.33 45.59 unst. unst. 735 0.52 0.93 3.71 7.60 10.47 16.04 30.39 unst. - - 858 0.46 0.88 2.70 9.29 15.20 unst. unst. - - - 981 0.47 1.13 4.39 unst. unst. - - - - - 1103 1.35 1.69 unst. ------1226 unst. unst. ------

Chapter 5: Bale Reactor Characterisation 204 5.4-1-4 Analysis and Calculation Procedure The RTD were acquired over adequate periods giving at least 10 mid-peak widths and complete return to initial conditions. Each RTD obtained required basic peak processing and the completion of the following tasks: 1. High frequency noise reduction (post processing: smoothing/filtering). 2. Threshold determination and baseline fitting, while the TCD had been dynamically zeroed generally a small deviation from zero needed correction for. 3. Cut off point determination for uniform analysis. 4. Basic moments calculation 5. Advanced method calculation and parameter tuning

Noise reduction in the time domain is commonly referred to as smoothing and that in the frequency domain as filtering. The exponential smoothing offered by Excel was immediately dismissed as it was observed that it shifted and distorted the data in time. Chow, Dave, Mulch and Yeow (1995), compared moving average and polynomial smoothing with rectangular and exponential low pass filters for signal enhancement of potentiometric stripping analysis using digital signal processing. Based upon their observations polynomial smoothing was chosen for further investigation as it had least effect upon area and gave little peak distortion retaining the characteristic features of the peaks.

Polynomial smoothing was popularised by Savitzky and Golay (1964). They tabulated the necessary convoluting integers in the form of smoothing tables. This operation can be mathematically represented as:

jm= Syj ()ij+ ysi() = ∑ (5.44) jm=− NORM

where NORM is a normalising factor, ys(i) is the value of the smoothed data point i, and

Sj are the tabulated values for a given window size w = (2m+1) and the order of the polynomial used to fit the portion of data. Application of this smoothing method thus results in the loss of m data points at the beginning and end of the data. Savitzky and

Golay tabulated Sj for w = 5 to 25. Values of NORM and Sj for the quadratic-cubic polynomial are given in Table 5.05. Using these a VBA program was written for

Chapter 5: Bale Reactor Characterisation 205 smoothing of the data once transferred from a text file generated by LabVIEW to Excel, Appendix A4.

Table 5.05 NORM and Sj values of Savitzky-Golay polynomial smoothing.

Points Sj ( w = 9 ) Sj ( w = 5 ) -4 -21 - -3 14 - -2 39 -3 -1 54 12 0 59 17 1 54 12 2 39 -3 3 14 - 4 -21 - NORM: 231 35

Through trial and error it was observed that: ™ repeated application of smoothing of a smaller window size, w = 9, retained more information and created less distortion than smoothing with a powerful large window size, w = 21. ™ application of smoothing beyond a certain number of times resulted in little further noise reduction and allowed the signal to converge to a given shape for that particular w.

Application of a 9-point filter 8 times was found to be optimum, over the range of gas and liquid flow rates studied.

Threshold determination and baseline smoothing were the first steps of the basic RTD analysis VBA program given in Appendix A4. After smoothing the data was prepared for further analysis and time was included. The column of time values was selected and became the reference column of the program. The program prompted for a region of data to be defined for which the average response was calculated and which became the threshold value. The rest of the data were zeroed according to this value. The method of moments was then applied.

In order to evaluate Per according to the method of Mecklenburgh and Hartland of Laplace transform of experimental data the following algorithm was applied to filtered and zeroed data:

Chapter 5: Bale Reactor Characterisation 206

1. Determine tm (method of moments):

∞ ∑ xoutt 0 tm = ∞ (5.45) ∑ xout 0

2. Transform the experimental data and determine gout choosing T as to just suppress the effect of tailing:

Tt ∞ − tm ∑ xoute 0 gout = ∞ (5.46) ∑ xout 0

3. Starting with the initial estimate of b = ln gout reiterate:

b2 Pe = (5.47) r bT+

4T a =+1 (5.48) P

⎡⎤⎛⎞a 2 ⎛⎞−+(1 aPe) r ⎢⎥41exp⎜⎟+−()a ⎜⎟ g 2 b =−ln ⎢⎥⎝⎠out ⎝⎠ (5.49) ⎢⎥()1+ a 2 ⎢⎥ ⎣⎦⎢⎥

until Per converges given a predefined tolerance ( abs (Pei - Pei-1) < 0.0001) .

Selection of T as to just allow for the suppression of tailing was carried out for Raschig and Bale packing gas phase data by comparing the amplitude of the transformed data to the average transformed response of the tailing portion of the peak (100 data points) and evaluating the ratio:

Chapter 5: Bale Reactor Characterisation 207 ⎛⎞n ⎜⎟∑ xout ⎜⎟in=−100 ⎜⎟100 ⎜⎟ α = ⎝⎠ (5.50) xout,max

Figure 5.09 (a) and (b) show the plot α, given by Equation (5.50), against T at various

ReG and ReL chosen across the full range of the experimental design. It can be seen that increased T suppresses the tailing portion of the peak to larger extent. T needs to be chosen such that suppression is just achieved as further increase in T erodes peak information. Based upon the studies value a value of T = 1.3 was adopted for Raschig ring runs and a value T = 1.6 was adopted for Bale packing runs.

Chapter 5: Bale Reactor Characterisation 208 0.03 ReG 8, ReL 245 0.025 ReG 56, ReL 490 ReG 20, ReL 981 0.02 ReG 129, ReL 0

α 0.015

0.01

0.005

0 0123 T (a)

0.007 ReG 8, ReL 245 0.006 ReG 56, ReL 490

0.005 ReG 20, ReL 981

0.004 ReG 129, ReL 0 α 0.003

0.002

0.001

0 00.511.52 T (b)

Figure 5.09 Measure of Gas Phase Tail Suppression α as a Function of the Laplace Parameter T for (a) Raschig and (b) Bale packing

Chapter 5: Bale Reactor Characterisation 209 5.4-2 Gas Phase Fluid Dynamics Results and Discussion The matrix design detailed in Section 5.4.1.3., with the limitation of flooding gave rise to 54 setpoints, which were repeated twice. The fundamental results are reported in Table 5.06 for Raschig Ring packing and Table 5.07 for Bale packing. Good reproducbility was observed for all RTD derived values.

Table 5.06 Basic Results of the Analysis of Gas Phase RTD curves for Raschig Ring Packing.

Set QG QL ReG ReL tm tm Per,G Per,G H0,G (L.min-1) (L.min-1) (s) (s) 1 0.28 0.00 8 0 239 247 17 13 0.48 2 0.28 0.50 8 245 285 269 11 12 0.55 3 0.28 1.00 8 490 250 265 11 10 0.51 4 0.28 1.50 8 735 274 261 10 12 0.53 5 0.28 2.00 8 981 262 267 8 9 0.52 6 0.28 2.50 8 1226 260 259 8 7 0.52 7 0.28 3.00 8 1471 236 241 7 7 0.47 8 0.66 0.00 20 0 110 110 23 23 0.51 9 0.66 0.50 20 245 115 116 16 16 0.53 10 0.66 1.00 20 490 115 114 15 15 0.53 11 0.66 1.50 20 735 116 117 13 13 0.54 12 0.66 2.00 20 981 107 107 12 13 0.49 13 0.66 2.50 20 1226 106 107 12 12 0.49 14 0.66 3.00 20 1471 105 107 10 10 0.49 15 1.88 0.00 56 0 52 52 45 45 0.69 16 1.88 0.50 56 245 54 54 20 19 0.71 17 1.88 1.00 56 490 54 55 22 18 0.72 18 1.88 1.50 56 735 54 53 16 17 0.70 19 1.88 2.00 56 981 52 52 14 14 0.68 20 1.88 2.50 56 1226 47 46 18 19 0.61 21 3.11 0.00 92 0 38 38 44 44 0.82 22 3.11 0.50 92 245 37 38 35 25 0.81 23 3.11 1.00 92 490 37 37 20 20 0.81 24 3.11 1.50 92 735 37 37 28 28 0.80 25 3.11 2.00 92 981 36 36 20 23 0.78 26 3.11 2.50 92 1226 35 35 22 22 0.76 27 4.33 0.00 129 0 30 30 37 37 0.91 28 4.33 0.50 129 245 29 29 29 31 0.88 29 4.33 1.00 129 490 29 28 28 30 0.86 30 4.33 1.50 129 735 28 26 26 34 0.83 31 4.33 2.00 129 981 28 26 20 29 0.83

Chapter 5: Bale Reactor Characterisation 210 Table 5.07 Basic Results of the Analysis of Gas Phase RTD curves for Bale Packing.

Set QG QL ReG ReL tm tm Per,G Per,G H0,G (L.min-1) (L.min-1) (s) (s) 32 0.28 0.00 8 0 245 246 23 22 0.49 33 0.28 0.50 8 245 251 214 8 9 0.46 34 0.28 1.00 8 490 234 246 7 7 0.48 35 0.28 1.50 8 735 240 243 6 7 0.48 36 0.28 2.00 8 981 234 231 6 6 0.46 37 0.66 0.00 20 0 95 97 29 28 0.44 38 0.66 0.50 20 245 87 87 10 10 0.40 39 0.66 1.00 20 490 85 84 10 10 0.39 40 0.66 1.50 20 735 84 84 10 10 0.39 41 0.66 2.00 20 981 81 77 9 10 0.36 42 1.88 0.00 56 0 49 50 36 34 0.66 43 1.88 0.50 56 245 44 44 18 19 0.58 44 1.88 1.00 56 490 44 44 14 13 0.57 45 1.88 1.50 56 735 41 41 15 16 0.54 46 1.88 2.00 56 981 41 42 16 14 0.55 47 3.11 0.00 92 0 38 38 27 28 0.83 48 3.11 0.50 92 245 35 35 17 15 0.76 49 3.11 1.00 92 490 33 33 14 14 0.72 50 3.11 1.50 92 735 33 33 14 13 0.71 51 4.33 0.00 129 0 25 25 34 34 0.76 52 4.33 0.50 129 245 23 23 21 20 0.70 53 4.33 1.00 129 490 23 23 18 17 0.71 54 4.33 1.50 129 735 21 21 17 16 0.63

With data acquisition rates of 12 to 25 samples per second the raw data was of too great a volume to be completely depicted for the complete matrix. Thus only a selection of RTD curves will be depicted and discussed. The selection was taken at the middle range of the study with reference values of ReG = 56 and ReL = 490. For ease of representation and file size the curves were sub-sampled giving a effective frequency of 1 sample/second.

For ReL 490 and various ReG filtered normalised RTD curves (E(t)) were reported in dimensionless time for Raschig and Bale packing Figures 5.10 (a) and (b) respectively. Qualitatively the peaks could be described as normal, with a degree of skewness and tailing. The tailing and observed shoulder could be an artefact of reactor stagnant zones. The normalisation process exaggerated the difference between the peaks, as the curves when captured were similar in shape. The difference in amplitude is not a dilution effect since higher flow rates gave more pronounced curves, but rather an indication of very

Chapter 5: Bale Reactor Characterisation 211 strong diffusion artefacts at low gas flowrates. The peaks were otherwise strong and singular with no indication of internal recirculation or significantly different parallel paths. The effect of ReG can be for both Raschig and Bale packing over this range of

ReG as one of promoting plug flow. The peaks became narrower and more upright giving a drop in variance and an increase in peakedness. The curves of Raschig and Bale packing were similar, with those of Bale marginally broader particularly at lower

ReG.

For ReG 56 and various ReL filtered normalised RTD curves (E(t)) were reported in dimensionless time for Raschig and Bale packing Figures 5.11 (a) and (b) respectively.

Increased ReL for both packing types led to quite obvious increased variance. And particularly for Raschig Ring packing increased tailing with a broader shoulder formed with increase in ReL. It thus seemed that liquid velocity promotes dispersion and backmixing and also creates stagnant spaces. These could be most easily visualised for Raschig Ring packing as an increase in the occurrence of small pockets of gas being trapped momentarily in the hollow centre of the Raschig Ring cylinder.

Chapter 5: Bale Reactor Characterisation 212 0.08

0.07 ReG 8 ReG 20 0.06 ReG 56 ReG 92 0.05

) ReG 129 m 0.04 E(t/t 0.03

0.02

0.01

0 00.511.522.53

t/tm (a)

0.09 ReG 8 0.08 ReG 20 0.07 ReG 56 0.06 ReG 92 )

m 0.05 ReG 129

E(t/t 0.04 0.03 0.02 0.01 0 0 0.5 1 1.5 2 2.5 3 (b) t/tm

Figure 5.10 Exit Age Distributions, E(t/tm), for various ReG at ReL = 490 of (a) raschig and (b) Bale packing

Chapter 5: Bale Reactor Characterisation 213 0.045 ReL 0 0.04 ReL 245 0.035 ReL 490 0.03 ReL 735 )

m 0.025 ReL 981

E(t/t 0.02 0.015 0.01 0.005 0 0 0.5 1 1.5 2 2.5 3 (a) t/tm

0.040 ReL 0 0.035 ReL 245 0.030 ReL 490 0.025 ReL 735 )

m ReL 981 0.020 E(t/t 0.015

0.010

0.005

0.000 01234(b) t/tm

Figure 5.11 Exit Age Distributions, E(t/tm), for various ReL at ReG = 56 of (a) raschig and (b) Bale packing

Chapter 5: Bale Reactor Characterisation 214

There is a more pronounced difference between E(t/tm) for ReL = 0 and ReL ≠ 0 for Raschig Ring packing than for Bale packing. While the packing was saturated before each run it is possible that Raschig rings drain much more and have a lower static holdup than Bale packing. Further the flow of gas may assist further drainage particularly in the case of Raschig rings. The static liquid of Bale packing is held between small spherical particles contained within fibreglass bags and hence wetness is better retained with the action of strong capillary forces and the open space does not vary significantly.

Given the near normality of the curves and the small deviation from plug flow the classical axial dispersion model was used to characterise the observed RTD. The dispersion type model can describe sufficiently the mixing behaviour in packed columns ranging from perfect mixing (Per → 0) to plug flow (Per → ∞). The suitability of this model has been demonstrated by many investigators, including: Marcias-Salinas and Fair (2000), Sater and Levenspiel (1966). Originally proposed by Danckwerts this model assumes that complete dispersion phenomena can be described by replacing the molecular diffusion coefficient by an effective dispersion coefficient. The latter accounts for the combined mixing effect caused by molecular and turbulent eddy diffusion, as well as by transfer of material in and out of the stagnant zones created by portions of fluid trapped inside the packing interstices, Marcias-Salinas et al (1999). The boundary conditions used here were those of a closed-closed system reflecting that dispersion occurs only along the length of the reactor and being the traditional choice for experiments conducted with external injection and sampling, Fogler (1999). The calculation procedure adopted was that promoted by Mecklenburgh and Hartland (1975). Which involves determining mean residence time by method of moments and

Per by Laplace transform of the data and solution against the known transfer function for pulse injection and external sampling.

Gas phase mean residence time, tm,G, varied as a function of both ReG and ReL, however as demonstrated by the bunching of the family of curves produced of Figure 5.12 (a) and (b) the effect of ReL was far less pronounced than that of the direct

Chapter 5: Bale Reactor Characterisation 215 300 ReL 0 ReL 245 250 ReL 490 ReL 735 200 ReL 981 ReL 1226

(s) 150 m t

100

50

0 0 50 100 150 (a) ReG

300 ReL 0 ReL 245 250 ReL 490 ReL 735 200 ReL 981

(s) 150 m t

100

50

0 0 50 100 150 (b)

ReG

Figure 5.12 Mean residence time as a function of ReG at different ReL for (a) Raschig Ring and (b) Bale packing

Chapter 5: Bale Reactor Characterisation 216 nonlinear influence of ReG of rapid decline in retention times. It is possible that with higher gas flowrates the gas increasingly penetrates a greater portion of the packing void space reducing the occurrence of stagnant pockets and creating a more uniform flow. It is also the action of these stagnant pockets to slow down the release of tracer at lower ReG hence giving longer tm.

For the same packed volume raschig rings displayed a marginally higher retention time for lower values of ReG (8 and 20) than Bale Packing. This was likely to be a consequence of the well-defined channels of wetted Bale packing and an indication that the particle portion of the packing is sealed to gas flow by the presence of the liquid. However at higher flowrates the retention times were very similar.

As mentioned the values of tm at ReL = 0 were those for a wetted bed. Efforts were made to wet the bed before each run. A condenser was used to prevent excessive liquid loss and prevent exposure of the GC to large amounts of moisture. A presaturator as typically used for mass transfer studies and as used for the basic capacity study was avoided in order to allow for injection as close to the entrance as possible and eliminate unnecessary end effects.

Gas phase total holdup, H0,G, was a derived value based upon tm,G. An overall reactor length of L = 1.5 m was used. Figure 5.13 (a) and (b) both suggest an “s” shaped curve in ReG and considerable variation in ReL. The gas holdup in general was higher for Raschig Ring packing than for Bale packing. The greater number of paths for the random packing compared to the well defined channels of the bale packing create higher retention.

Chapter 5: Bale Reactor Characterisation 217 1 ReL 0 ReL 245 0.9 ReL 490 ReL 735 0.8 ReL 981 ReL 1226

0,G 0.7 H

0.6

0.5

0.4 0 50 100 150 (a)

ReG

1

0.9 ReL 0 0.8 ReL 245 ReL 490 0.7 ReL 735 0.6 ReL 981

0,G 0.5 H 0.4 0.3 0.2 0.1 0 0 50 100 150 (b)

ReG

Figure 5.13 Gas Phase Total Holdup as a function of ReG at different ReL for (a) Raschig Ring and (b) Bale packing

Chapter 5: Bale Reactor Characterisation 218 Holdup was expressed in terms of ReG and ReL as a multivariable linear expression:

−−35 H0,GRR , =+×0.507 3.123 10 ReG −× 3.732 10 ReL (5.51) and

−−35 H0,GB , =+×0.461 2.570 10 ReG −× 8.869 10 ReL (5.52)

The use of gas phase holdup or porosity is required for the calculation of interstitial packing velocity.

The reactor Peclet number, Per, was used as the fundamental measure of mixing within this study. It allowed for comparison of the bulk effects of the packing choice. Figure

5.14 (a) and (b) showed the variation of Per,G with ReG for different ReL. These plots showed that the reactor behaves as a plugflow vessel (Pe → ∞) with increase in gas flow irrespective of the liquid flowrate. Bale packing, however, appeared to improve axial gas phase mixing. Although to lesser extent reactor Peclet number, Per, decreased almost linearly with increased ReL as evidenced from Figure 5.15 (a) and (b). This behaviour is consistent with what one would expect in the preloading regime as the higher liquid flow rate forces gas phase backmixing, as it displaces gas from the void space.

Chapter 5: Bale Reactor Characterisation 219 35

30 ReL 245 25 ReL 490 ReL 735 20 ReL 981 r,G

Pe 15

10

5

0 (a) 0 50 100 150

ReG

21

19 ReL 245 17 ReL 490 15 ReL 735

r,G 13 Pe 11

9

7 5 (b) 0 50 100 150

ReG

Figure 5.14 Gas Phase Reactor Peclet number, Per,G, as function of ReG for

different ReL for (a) Raschig Ring and (b) Bale packing.

Chapter 5: Bale Reactor Characterisation 220 50 45 ReG 8 ReG 20 40 ReG 56 35 ReG 92 30 ReG 129

r,G 25 Pe 20 15 10 5 0 (a) 0 500 1000 1500

ReL

40 ReG 8 35 ReG 20 30 ReG 56 ReG 92 25 ReG 129

r,G 20 Pe 15

10

5

0 (b) 0 200 400 600 800 1000

ReL

Figure 5.15 Gas Phase Reactor Peclet number, Per,G, replotted as a function of ReL

for different ReG for (a) Raschig Ring and (b) Bale packing.

Chapter 5: Bale Reactor Characterisation 221 For closed vessels Levenspiel (1972) proposed an indicative scale for Per as reported in Table 5.08.

Table 5.08 Indication of extent of backmixing, Levenspiel (1972)

Per Implication ∞ plug flow 500 small amount of dispersion 40 intermediate amount of dispersion 1 large amount of dispersion 0 mixed flow

In light of Table 5.08, Raschig Rings and Bale packing over this range of ReG and ReL were found to give intermediate to large amounts of backmixing. The majority of observed Per fell in the range 10 to 30 indicating moderate amounts of dispersion.

Consistent with previous qualitative observations, Per,G values from the first generation packing were generally higher then those for Bale packing. Even so, within each packed bed, Per.G seemed to increase with increased gas flow rate but decreased with increased liquid flowrate. For a fixed liquid velocity, it is conceivable that gas backmixing will occur at lower gas flowrates as the upward moving gas struggles to find its way through the unwetted void spaces in the bed. As a result, the transport response curve is characterised by a long tail. However, as the ReG increased, the gas momentum became sufficient enough to ensure a near plug flow situation through the void region of the bed as evidenced by the narrower but longer RTD curves. By the same token, as liquid downward flowrates increases, one would expect greater axial mixing in the upward moving gas phase and this was indeed observed.

Due to the lack of fundamental theory for axial dispersion in countercurrently operated packed beds the results were correlated by means of equations derived from dimensionless analysis. Commonly proposed for Per,G correlation within countercurrent packed bed systems is equation given in general form as:

δ β χ ReL Pe = α ReGp .10 .(d ap) (5.53)

Chapter 5: Bale Reactor Characterisation 222 Equation (5.53) suggests power law behaviour in ReG as observed here and allows for the incorporation of ReL = 0 to be used, while capturing the parallel gradually stepping family of curves resulting from variation of ReL. The term dpap is used to account for packing geometry of random packing. Similar terms have been developed for structured packing which incorporate angle of inclination and mesh geometry. It should be noted at this point that Bale packing as yet has not been standardised. The design, materials and tightness of wrapping are highly variable thus each system needs individual characterisation. This is a drawback for Bale packing. There are no clear well-defined variables allowing for control of holdup by packing manipulation for current CD hardware designs. In structured packing such as KATAPAK-S there is the channel angle relative to the vertical axis, but even this gives only weak design control Moritz and Hasse (1999). As dpap was not varied explicitly for this study it was not used as a correlating variable.

Consequently the combined effect of both gas and liquid flow dynamics upon the mixing characteristic Per,G were adequately captured by:

0.35 −0.00026ReL PerGRR,, = 7.69ReG .10 (5.54) and

0.19 −0.00055ReL PerGB,, =13.04ReG .10 (5.55)

Both Equation (5.54) and (5.55) gave adequate experimental data representation as evident from the parity plots of Figure 5.16 (a) and (b). The greater variance of Equation (5.55) for Bale packing can be associated with the multiple levels of porosity of Bale packing which depending on random events may lead to very different flow paths. Although the gas flow through the Bale packing will occur predominantly through the open channels created by the rolled quilt and mesh, with lower levels of wetting the gas can penetrate across the fibreglass bag and experience flow through tightly packed spherical particles. Hence, while Raschig Ring packing is packed randomly a certain degree of uniformity is created.

Chapter 5: Bale Reactor Characterisation 223 50

40

30 Model r,G 20 Pe

10

0 0 1020304050 (a) Per,G Experimental

40

30

Model 20 r,G Pe

10

0 0 10203040

Per,G Experimental (b)

Figure 5.16 Parity plots of Equations (5.54) and (5.55) modelling Gas Phase

Reactor Peclet number, Per,G, for (a) Raschig Ring and (b) Bale packing.

Chapter 5: Bale Reactor Characterisation 224 Equation (5.54) for Raschig Ring packing was also reformulated for Peu allowing for comparison to published correlations for gas phase dispersion.

0.14 −0.00024ReL PeuGRR,, = 0.13ReG .10 (5.56)

In this form Sater and Levenspiel for dp 12 mm obtained:

−0.79 −0.00299ReL PeuGRR,, = 6.23ReG .10 (5.57)

over the range 51 ≤ ReG ≤ 180 and 0 ≤ ReL ≤ 300. They observed increased mixing for an increase in ReG and increased mixing for an increase in ReL. For very high ReB,G

(340 ≤ ReB,G ≤ 4066) and dp 25 mm Marcias-Salinas and Fair (2000) also observed and correlated increased mixing with ReG. For both packings under current conditions decreased mixing was observed as ReG was increased indicating a tendency towards plug flow. Increased mixing was observed as ReL was increased. Thus, while the liquid velocity contribution behaves in a similar manner, the gas velocity gave a reverse trend. It should be noted, that in terms of absolute values for the current study interstitial packing based Peclet number, Peu, ranged from 0.07 to 0.3, while all the values of both Sater and Levenspiel and Marcias-Salinas and Fair were above unity and displayed little variation. Being above unity they could be considered to be very close to plug flow and only displaying small amounts of dispersion. Thus, the assumption of plug flow for higher vapour velocities found with CD operation in the pre-loading regime may be adequate.

Dispersion is affected by both molecular and eddy diffusion. For single gas phase flow through granular beds it was shown by Edwards and Richards (1968) that for low

Bodenstein type Reynolds numbers, ReB,G < 1, the dispersion is only affected by molecular diffusion, but at higher values, ReB,G > 10, eddy diffusion predominates. Marcias-Salinas and Fair (2000) proposed that for the conditions of their investigation eddy diffusion would completely predominate and hence the increase in turbulence and greater dispersion at higher gas velocities was observed. They suggested that the

Chapter 5: Bale Reactor Characterisation 225 increase of mixing at high gas flow rates could be explained by the formation of more eddies in the gas flow channels through the packing.

An extension of this argument would suggest that under the current conditions for both types of packing molecular diffusion predominated. That as velocity is increased there is a tendency towards plug flow over this range of ReG. The range of ReB,G covered was 1 to 23 for 8 mm Raschig Ring packing. The openness of Bale packing (ε = 0.7171) and 8 mm Raschig Rings (ε = 0.7257) may result in near continuous channels being established, disturbed only by high flowrates of liquid. However, the curvature of the

Per curves in ReG indicates that eddy diffusion is starting to influence the mixing behaviour. Further that over a wider range of ReG an eddy diffusion dominated region may be reached where the trend would be reversed.

A wider investigation may show that indeed across the full span of ReG there are various dispersion regimes for gas phase flow in the presence of liquid phase flow for Raschig Ring packing and Bale packing. It is however difficult to find an experimental method, which will span a wide range of ReG or ReL. For the current study although interested in predominantly low to moderate velocities associated with CFBR operation, the injection and detection methods were limited to ReG 130. In the case of Marcias- Salinas and Fair (2000) the membranes used to sample within the packing studied failed at high liquid and low gas velocities. Recently FTIR probes have been developed which allow in-situ reactor IR analysis through the use of advanced optics contained with a d =

10 mm probe. Their application for non-aqueous systems using CO2 as a tracer may offer a wider range of velocities and liquid phase properties.

Chapter 5: Bale Reactor Characterisation 226 5.5 Liquid Phase Mixing in the Pre-Loading Regime Liquid phase mixing was study in the preloading regime using a newly developed method involving digital photography as an extension of the standard spectrophotometric techniques.

5.5-1 Experimental The experimental method used was developed from scratch to allow for the determination of residence time distribution and holdup in the reactor available with the packing of choice Bale packing and Raschig Ring packing. Again pulse injection at inlet and sampling at exit were adopted as the experimental approach. The use of two- point analysis was considered, however sampling of the bale packing proved difficult. Internal sampling of the reactor was seen as necessary as the liquid collector at the bottom of the countercurrent system would significantly alter the overall exit distribution of the column.

It was noticed that the existing reactor packing support, consisting of four centrally pointing arms, channelled the down flowing liquid towards the wall of a small uniform window like section of the glass reactor located immediately before the bottom quick fit ground glass connection. It was recognised that this window provided a unique sampling opportunity, which could be utilised with some form a spectrophotometric approach. The following methods were considered: ™ laser absorption ™ digital video and microscopy frame analysis ™ digital photography and microscopy photo analysis

Laser absorption although cheap and continuous was discounted due to the relatively poor optical properties of the sampling window. Digital video was investigated, however at the time adequate technology allowing for frame by frame analysis of images of high resolution was more expensive than could be justified for the project. During the investigation of scientific video image processing, software such as Image-J and ScionImage were acquired, explored and relevant examples worked through.

Chapter 5: Bale Reactor Characterisation 227 Conventional digital photography was considered and it was found that adequate sampling rates could be obtained.

5.5-1-1 Liquid Phase Procedure The rig developed was depicted in Figure 5.17. The experiments were conducted using water and air at ambient temperature and atmospheric pressure. Potassium permanganate dissolved in water was chosen as tracer, due to availablity, previous laboratory experience and strong broad adsorption of green light. The absorption of green light from a tungsten light source was measured by taking photos of a convenient sample window at the base of the reactor using a digital camera and analysing these using the microspectroscopy software ScionImage.

This method allowed for rapid rates of sampling. Photos were taken in a rapid burst mode with 16 photos taken every 5 seconds. By repeating the experiment an overall RTD response was obtained. The overlap between sampling periods proved to be excellent. In order to allow for a relative comparison between photos and a constant base line the camera was set to manual mode. The manual mode here is a virtual one and its settings were listed in Table 5.09. The photos were captured at 0.8 Mega Pixels in JPEG format and converted to TIFF format using the software Irfan View, which was capable of batch conversion. Electronic noise was removed from the raw data with a 5- point Savitzky-Golay filter.

Table 5.09 Kodak DC4800Camera settings for liquid RTD determination Setting Value/Selection Aperture, f 2.8 White Balance - Quality saturated Resolution 0.8 Mega-Pixel Exposure Distribution Centre Weighted Sharpness Sharp Film Sensitivity ISO 200 Exposure 1/125 Date none

Chapter 5: Bale Reactor Characterisation 228 vent

[04]

[01] [05]

[02]

[06]

[03] [09]

R [07] [08] Water Tank N2 [10] Drain

Figure 5.17 Rig for Experimental Determination of CFBR Liquid Side Fluid Dynamics

Key to Figure 5.17 CFBR Liquid Side Fluid Dynamics [01] Liquid Flow Meter and Controller, BASIS Micro-Motion M/N F025S1319SM and Eurotherm 2216 PID controller. [02] Air actuated control valve and current to pressure transducer, Fisher Controls, M/N 912346. [03] Centrifugal pump, Grundfos, CH12-40-A-W-6-BQQV, 640 kW, h 37m. [04] Injection port with septum. [05] Trickle Bed Reactor jacketed column. [06] Dark space enclosure. [07] Tungsten light source, 25 W. [08] Diffuser, a translucent mat plastic sheet. [09] Digital Camera, Kodak DC4800.

Chapter 5: Bale Reactor Characterisation 229 [10] Nitrogen gas supply and rotameter.

Liquid Side Fluid Dynamics Stepwise Operating Procedure 1. Set gas flow rate as required using the gas side rotameter [10]. 2. Set the liquid flowrate using the controller [01], allow time for steady state to be achieved. 3. Prepare the camera making sure the memory card is free, the camera is set to manual and the correct settings have been entered and that the power source is connected. 4. Inject 5 mL of 0.2 M potassium permanganate and initialise the camera’s burst mode simultaneously. 5. Using a stopwatch take photos of the next 5 sec interval and repeat this procedure until a complete curve is gained. Make sure to leave adequate time for the tracer to clear the reactor before commencing a further interval. 6. Download the photos in an established directory and file. 7. Convert the images from JPEG to TIFF using InfranView and it’s batch processor. 8. Open a sample image in ScionImage and define the sample window in terms of pixel coordinates of the photo and width and height (X,Y,W,H). Enter these settings into the Batch Processor program written (Appendix A3) and calculate the green light intensity. A text file should be generated containing the results as a single column of numbers. 9. Import the text file data into Excel, average every 16th entry and run the Savitzky- Golay filter program (Appendix A4). 10. Next run the Baseline Adjustment and Method of Moments programs (Appendix A4). 11. Sub-sample the filtered data and feed this to Polymath 5 for nonlinear regression against the nominated fluid dynamics model

Chapter 5: Bale Reactor Characterisation 230 5.5-1-2 Experimental Design Both Raschig Ring and Bale packing were investigated for the same reasons as outlined in Section 5.4.. The range of Reynolds number investigated was the same as for the gas side fluid dynamics discussed in Section 5.4.1.3.. However, given that literature suggested that tm, H0,L and Per did not vary in ReG for the pre-loading regime a complete matrix design was not considered necessary. Instead the this assumption was qualified at ReL = 735 for various ReG. Further runs were conducted at ReG = 0 for:

ReL : 0, 245, 490, 735, 981, 1226 and 1471

5.5-1-3 Calculation Procedure The digital images were treated in ScionImage (processing code Appendix A3). Figure 5.18 demonstrates a sequence of images and associated raw absorption values. The calculation procedure for the liquid phase was much the same as that for the gas phase described in Section 5.4-1-4. However given the lower volume of data a Savitzky-Golay filter of window size w = 5 was only applied once (Appendix A4). The values were filtered, the baseline was zeroed and an E(t) curve was generated. The effect of the filter is illustrated in Figure 5.19 for ReL =735 and ReG = 0. Further the Laplace transform was applied with a moderate value of T = 0.5 giving a very small yet adequate amount of tail suppression.

Chapter 5: Bale Reactor Characterisation 231 t=0, y=41.69 t=5 s, y=99.56

t=10 s, y=118.49 t=15 s, y=85.20

t=20 s, y=61.91 t=30 s, y=45.97

Figure 5.18 Example Images and Values obtained, Bale packing ReL =735

Chapter 5: Bale Reactor Characterisation 232 160 Raw Data 140 Filterred Data 120 Baseline Corrected

100

80

60 Signal (gray scale) 40

20

0 0 10203040 t (s)

Figure 5.19 Calculation example corresponding to

images of Figure 5.18, Bale packing with ReL = 735

Chapter 5: Bale Reactor Characterisation 233 5.5-2 Liquid Phase Fluid Dynamics Results and Discussion The design detailed in Section 5.5.1.3., with the limitation of flooding gave rise to setpoints as presented. The results are reported in Table 5.10 for Raschig Ring packing and Table 5.10 for Bale packing. The variation of tm, H0,L and Per,L with independent parameters ReG and ReL will be discussed in the subsequent section.

Table 5.10 Basic Results of the Analysis of Liquid Phase RTD curves for Raschig Ring Packing.

Set QG QL ReG ReL tm Per,L H0,L (L.min-1) (L.min-1) (s) 1 1.9 1.5 56 738 10.05 34.10 - 2 3.1 1.5 92 738 10.12 33.53 - 3 4.3 1.5 129 738 10.36 27.37 - 4 5.6 1.5 165 738 9.88 29.03 - 5 6.8 1.5 201 738 10.33 32.83 - 6 8.0 1.5 238 738 10.30 29.49 - 7 9.2 1.5 274 738 9.72 31.41 - 8 0 0.5 0 246 16.81 41.25 0.15 9 0 0.75 0 369 13.39 15.69 0.18 10 0 1 0 492 11.43 19.90 0.20 11 0 1.25 0 615 10.94 24.46 0.24 12 0 1.5 0 738 10.07 29.96 0.26 13 0 1.75 0 861 9.45 33.15 0.29 14 0 2 0 983 8.73 37.29 0.30 15 0 2.25 0 1106 8.56 40.59 0.34 16 0 2.5 0 1229 8.04 43.60 0.35 17 0 2.75 0 1352 7.89 48.02 0.38 18 0 3 0 1475 7.58 50.87 0.40

Chapter 5: Bale Reactor Characterisation 234 Table 5.11 Basic Results of the Analysis of Liquid Phase RTD curves for Bale Packing

Set QG QL ReG ReL tm Per,L H0,L (L.min-1) (L.min-1) (s) 19 0.28 1.50 8 738 10.38 8.46 - 20 0.66 1.50 20 738 12.05 6.54 - 21 1.88 1.50 56 738 12.25 8.03 - 22 3.11 1.50 92 738 12.28 7.86 - 23 4.33 1.50 129 738 12.12 7.71 - 24 0.00 0.50 0 246 28.81 15.31 0.05 25 0.00 0.75 0 369 21.01 12.95 0.07 26 0.00 1.00 0 492 16.76 9.28 0.09 27 0.00 1.25 0 615 15.19 6.44 0.13 28 0.00 1.50 0 738 11.52 6.95 0.10 29 0.00 1.75 0 861 12.13 7.35 0.17 30 0.00 2.00 0 983 10.14 9.17 0.15 31 0.00 2.25 0 1106 9.46 9.55 0.17

The lack of dependence of the tm and Per upon ReG in the pre-loading regime was qualified in set 1 to 7 and 19 to 23 for Raschig and Bale packings respectively. The mean, standard deviation, 95 % confidence interval and R2 of a linear regression were reported in Table 5.12. It can be seen that a tight distribution affected by experimental error to a greater extent than variation in ReG for each variable is formed, with very low R2 values. Figure 5.20 (a) and (b) also demonstrated graphically the lack of dependence of Per upon ReG. Per was considered an important indicator as it was considered that

ReG may influence mixing without affecting tm to an appreciable extent.

Table 5.12 Lack of variation of dependent variables in ReG Packing: Raschig Ring Bale Packing Variable: tm Per tm Per mean 10.10 30.96 11.77 7.59 stdev σ 0.23 2.39 0.73 0.71 ±k 95% 0.16 1.65 0.59 0.57 R2 0.0185 0.0206 0.0882 0.0787

Chapter 5: Bale Reactor Characterisation 235 40

35

30

25

r.L 20 Pe 15

10

5

0 0 100 200 300

ReG (a)

9 8 7 6 5 r.L

Pe 4 3 2 1 0 0 50 100 150

ReG (b)

Figure 5.20 Liquid Phase lack of dependence of Per,L upon ReG for (a) Raschig Ring and (b) Bale packing

Chapter 5: Bale Reactor Characterisation 236 3 ReL 245 2.5 ReL 490 ReL 735 2 ReL 981 )

m ReL 1226 1.5 ReL 1471 E(t/t 1

0.5

0 0 0.5 1 1.5 2 2.5

t/tm (a)

1.2 ReL 245 1 ReL 368 ReL 490 0.8 ReL 613 ) m 0.6 ReL 735

E(t/t ReL 981 0.4 ReL 1103

0.2

0 01234

t/tm

Figure 5.20 Liquid Phase Exit Age Distributions as a Function of ReL for (a) Raschig Ring and (b) Bale packing.

Chapter 5: Bale Reactor Characterisation 237 The newly developed method based upon digital photography proved successful for the range of flowrates of interest. The exit age distributions, E(t/tm), obtained through analysis of the RTD acquired as a function of ReL are presented in Figure 5.21 (a) and (b) for Raschig and Bale packing respectively. Extensive tailing, multi peaks or other common curve traits suggesting reactor non-idealities were not apparent. The characteristics of observed curves suggest applicability of the axial dispersion model.

Raschig ring E(t/tm) approached plug flow with increase in ReL, with highly symmetrical distributions evolving centred about unity. Bale packing showed a greater deviation from plug flow and considerable skewness as well as a greater degree of mixing. It was difficult to determine a trend in ReL through mere inspection of E(t/tm).

Liquid phase mean residence time, tm,L, decreased non-linearly in ReL for both types of packing as may be seen in in Figure 5.22 (a) and (b). The retention of liquid by Bale packing compared to 8 mm Raschig rings was marginally higher, however the overall behaviour was very similar. The porosity (voidage) of Bale packing is 0.7171 while that of Raschig ring packing is 0.7257.

The total liquid holdup, H0L of the bed defined by Equation (5.17) was calculated based upon the mean residence time obtained from the RTD data. Figure 5.23 (a) and (b) show the variation of H0L with ReL in the Raschig ring and Bale packing respectively.

The power law behaviour of HD,L in ReL in first generation randomly packed beds has also been observed previously by many investigators thus:

m HH0,LSLLL=+ , α Re (5.58)

where HS,L was the static liquid holdup, namely, the volume fraction of resident stagnant liquid within the packing interstitial space due to capillary forces. As a result m HS,L is only a function of the packing characteristics. The second term αLReL is the dynamic liquid holdup contribution, HD,L, which determines the mixing mechanism,

Chapter 5: Bale Reactor Characterisation 238 16

14

12

10

(s) 8 m t 6

4

2

0 0 500 1000 1500 2000

ReL (a)

35

30

25

20 (s) m t 15

10

5

0 0 500 1000 1500

ReL (b)

Figure 5.22 Liquid Phase tm as a function of ReL for (a) Raschig Ring and (b) Bale packing.

Chapter 5: Bale Reactor Characterisation 239 0.35

H = 7.692U 0.9226 0.3 D,L L

0.25

0.2 D,L H 0.15

0.1

0.05

0 0 0.01 0.02 0.03 0.04 (a) UL

0.18 245 < ReL < 1471 0.16 15 < ReL < 100 0.14 0.12

L 0.1 , D

H 0.08 0.06 0.04 0.02 0 0 0.005 0.01 0.015 0.02 0.025

UL (b)

Figure 5.23 Liquid Phase Dynamic Holdup, HD,L as a function of UL for (a) Raschig Ring and (b) Bale packing.

Chapter 5: Bale Reactor Characterisation 240 molecular diffusion, viscous effects or eddy diffusion. The following correlations were developed for H0L for Raschig and Bale packing:

−04 0.92 H0LRaschig =+×0.09 4 10 ReL (5.59) and

−04 1.02 H0LBale =+×0.22 1 10 ReL (5.60)

As Bale packing was fabricated of fibreglass bags containing 850 μm glass spheres with a high potential for liquid retainment, HSL estimated, as 0.22 is higher than that in

Raschig at 0.09. Interestingly αL for Bale packing is about one quarter of that in Raschig rings packing, which supports film and rivulet flow along the downward channels as a major mode of flow. Liquid hold-up values are also comparable to those found by other investigators. The value for Bale packing static holdup of 0.22 also compares well with that of instant and rested static holdup: 0.2203 and 0.2047 determined gravimetrically in section 5.2-2 for Bales outside of the reactor.

Compared to the basic capacity study conducted at low ReL (15 to 100) the liquid retention was much fuller and almost linear with liquid velocity. The previous data are also plotted in Figure 5.23 (b) for comparison. The values being very close to that of static holdup and determined volumetrically would give large errors if they were extrapolated beyond the range of ReL of the basic capacity study.

The reactor Peclet numbers, Per,L, were plotted as a function of ReL, Figure 5.24 (a) and

(b). Over the Raschig ring bed, Per,L increased monotonically with ReL as evidenced from Figure 5.24 (a). The values followed a power law type trend with a minor yet significant extent of curvature.

Chapter 5: Bale Reactor Characterisation 241 60 0.8538 PeL = 0.1021ReL 50 R2 = 0.9974 40 L 30 Pe

20

10

0 0 500 1000 1500 (a) ReL

18 16 14 12 10 r,L

Pe 8 6 4 Region 1 Region 2 ReL > 623 2 ReL< 623 0 0 250 500 750 1000 1250 (b) ReL

Figure 5.24 Liquid Phase Reactor Peclet number, Per,L, variation in ReL for (a) Raschig Ring and (b) Bale packing.

Chapter 5: Bale Reactor Characterisation 242 Thus a power-law correlation was used to describe PeL for Raschig rings given by:

0.85 PerLR,, = 0.102ReL (5.61)

The adequacy of the correlation is evident in Figure 5.24 (a).

As detailed in the background theory (Section 5.2) there a number of ways of expressing Peclet number and the associated independent dimensionless variables of gas and liquid Reynolds number. Their form creates ease of use in final application and normalises the effects involved allowing for better comparison of different packing types of the same overall series. In order to allow for comparison with published data

Equation (5.61) was expressed in the interstitial Peu hence:

−03 0.85 PeuLRR,, =×1.702 10 ReL (5.62)

The experimental values of Peu,L,RR and Equation (5.62) were compared against published correlations in Figure 5.25. These correlations and their range were reported in Table 2 of the literature review. They were obtained over Raschig rings of various sizes under similar conditions and are generally appropriate over the range of ReL of comparison. Where correlations were given in Reu,L they were converted to ReL through the use of reported Raschig ring size and column diameter. It became immediately evident that the correlations in the ReL form were highly dependent upon packing nominal size. The smaller packing sizes gave increased interstitial mixing as may be expected when considering ReL. The smaller packing creates a greater possibility of fluid paths leading to enhanced mixing. Additionally, the smaller packing occupies more space giving a lower porosity and less open space allowing for the establishment of continuous channels of flow. The correlation of Stockar et al (1984) followed obviously different power law behaviour than the other correlations. Their data was obtained using organic solvents and the difference in behaviour of these to aqueous systems may be poorly accounted for by Re and Ga.

Chapter 5: Bale Reactor Characterisation 243 2.5

2

1.5 u,L

Pe 1

0.5

0 300 800 1300

ReL

Present Work 8 mm Furzer (1972) 6.4 mm Otake (1958) 7.6 mm Bennett (1970) 6.4 mm Sater (1966) 12 mm Stockar (1984) 25 mm Macias-Salinas (2000) 25 mm

Figure 5.25 Raschig Ring Fluid or Packing Peclet number as a function of Reu,L and comparison with: Furzer (1972), Otake (1958), Bennett (1970), Sater (1966), Stockar (1984), and Macias-Salinas (2000).

Chapter 5: Bale Reactor Characterisation 244 Based upon these comparisons it was concluded that the newly developed method gave results of similar accuracy to previously employed published methods, such as traditional and flow-cell spectrophotometry and conductivity measurements and formed a continuity with previously published Raschig Ring data.

The comparison revealed that a unifying umbrella correlation might be developed to account for packing size difference. For the one series or type of packing varying only in size it was found that the commonly used geometry variable (apdp) showed little variation and could not easily account for the observed difference in Peu,L. Instead the size dp = 25 mm = 1” was defined as a reference size, dp,ref. It is typically as large as pilot scale operation allows for. While industrial applications including absorption and distillation use packing in the range 13 to 76 mm (0.5” to 3”) it is doubtful whether CFBR application would use catalytic Raschig Rings of size greater than 13 mm given the requirements of internal mass transfer. Using this aspect ratio the following correlation was proposed for operation with Raschig rings in the pre-loading regime:

1.179 ⎛⎞d Pe = 0.0549Re0.49 p uLRR,, L ⎜⎟ (5.63) ⎝⎠d pref,

2 where dp,ref = 25 mm and R = 0.90.

While Per,L increased almost linearly with ReL for the Raschig ring bed, the behaviour in the Bale packing suggests two distinct mixing regimes. As shown in Figure 5.24, at low liquid flowrates ( 250 < ReL < 623) axial dispersion increased with increasing liquid velocity. However beyond a critical ReL of about 623, Per,L rose almost linearly with increasing ReL indicating a significant tendency towards liquid plug flow in the Bale packing beyond this point. The degree of dispersion resembled that at the lower ReL of 8mm Raschig rings. Given the comparison of Figure 5.25, 6.4 mm (1/4“) Raschig rings may give a similar degree of dispersion. As evidenced by Table 5.08 the range of dispersion may be regarded as intermediate to large.

Due to the more open structure of the Bale packing at very low velocities, liquid simply flowed down the Bale bed as rivulets along the line of least resistance, with almost plug

Chapter 5: Bale Reactor Characterisation 245 flow conditions. This occurred centrally, as wall wipers were installed three per bale, to prevent short cutting along the wall. As the liquid flow rate was increased the fraction of bed void space occupied by the liquid would increase, causing greater mass exchange between the stagnant liquid pools within the bags and the main liquid flow.

Conceivably, at a critical ReL, maximum mass transfer between the mobile liquid phase and the stagnant isolated pools would be attained. This is the point at which Per,L is at its minimum. As the liquid flow increases beyond this point, mixing via molecular transport became less important and viscous effects became more dominant as the bulk of the liquid was now in the mobile phase. Thus the liquid entering the column can flow within the packed channels of the bags with longer residence times or the open channels with shorter residence time and there is exchange between the two modes of flow. It can be suggested that operation in Region 2 at ReL just beyond ReL = 623 would give optimum exchange.

For the purpose of correlation it was desirable to retain obvious information about the two regions of mixing and thus a linear fit was assigned to each giving:

Region 1. () 250≤≤ ReL, 623 : PerLB, = 21.59 − 0.025ReL (5.64) and

−03 Region 2. () 623 ≤≤ ReL,, 1226 : PerLR = 1.9738 +× 6.9 10 ReL (5.65)

Theses regions of Per,L,B for 43×100 mm Bale packing are important for relatively short beds, such as the current bed of L = 0.6 m and for most practical lengths. At the minimum Per,L corresponding to ReL = 623, the value of Per,L will only reach the intermediate to small range of dispersion at a bed length of 12 m. Only well beyond this length will the significance of the regions be diminished and plug flow may be assumed.

5.6 Concluding Remarks It was found that in comparison to first generation random packing of nearly equal porosity, Bale packing displayed higher amounts of liquid retention and backmixing. It also gave higher amounts of gas phase backmixing. This behaviour could be attributed

Chapter 5: Bale Reactor Characterisation 246 to the multiple levels of porosity of Bale packing, in particular well defined channels and many channel structures filled with tightly packed small particles.

Loading and flooding for Bale packing over a range corresponding to distillation operation of 500 ≤ ReG ≤ 3000 and 10 ≤ ReL ≤ 100 were defined by:

−0.78 ReG, Loading = 21580ReL (5.41) and

−0.35 ReG, Flooding = 7357 ReL (5.42)

Loading and flooding for relatively high liquid flowrates and low gas flowrates over the range 188 ≤ ReG ≤ 456 and 0 ≤ ReL ≤ 1226 occurred simultaneously and was defined by:

−1.109 ReGB..= 580633ReLB (5.43)

Correlations were developed for holdup and gas and liquid phase reactor Peclet number for operation of Bale packing within the pre-loading regime. The liquid phase holdup of Bale packing in the preloading regime was correlated by:

−04 1.02 H0LBale =+×0.22 1 10 ReL (5.59)

For a 0.6 m bed of 43×100 mm Bale packing gas phase backmixing was well described by:

0.19 −0.00055ReL PerGB,, =13.04ReG .10 (5.54)

which returned a positive exponent in ReG suggesting approach to plug flow with increased ReG and suggested operation under diffusion predominated conditions for the range of investigation.

Chapter 5: Bale Reactor Characterisation 247

The liquid phase backmixing in the preloading regime displayed a clear minimum, which defined two regions, and these were well described by:

Region 1. () 250≤≤ ReL, 623 : PerLB, = 21.59 − 0.025ReL (5.63) and

−03 Region 2. () 623 ≤≤ ReL,, 1226 : PerLR = 1.9738 +× 6.9 10 ReL (5.64)

This behaviour was considered to be a symptom of multiple levels of porosity of bale packing and the exchange between film flow along bag walls and packed flow.

Correlations for holdup and backmixing were employed for reactor analysis of CFBR and CED experimental runs of Chapter 7. Correlations for loading and flooding were employed when determining the regime of operation for the extractive distillation study considered in the next chapter, Chapter 6.

Chapter 5: Bale Reactor Characterisation 248 5.7 Nomenclature

2 -3 ap packing surface area (m .m ) A, B constants defined in correlations

Bo Bodenstein number, vL/Da C tracer concentration, (mol.L-1) -1 C0 tracer injection concentration, (mol.L ) d reactor diameter, (m) dp particle diameter, nominal packing diameter, (m)

Da axial dispersion E(t) exit age distribution 3 2 -2 Ga Galileo number (dp .g.ρ .μ ) -1 H0 total holdup (vol.vol ) -1 HD dynamic holdup (vol.vol ) -1 HS static holdup (vol.vol ) I(t) internal age distribution L bed length, (m) m, n exponents defined in correlations P pressure (kPa) -1 Per reactor Peclet number (= UL.Da ) -1 Peu particle Peclet number (= Vdp.Da ) Re superficial velocity reactor Reynolds number (= ρUd.μ-1) -1 Reu particle Reynolds number (= ρVdp.μ ) t time (s) tm mean residence time (s) T Laplace Parameter U superficial fluid velocity, (m.s-1) -1 v interstitial fluid velocity (=L/tm), (m.s ) x, y, z coordinates

Subscripts G gas phase L liquid phase

Chapter 5: Bale Reactor Characterisation 249 Greek Letters α tail suppression parameter λ(t) intensity function

λm mean residence intensity μ viscosity, Ns.m-2

θ dimensionless time (t/tm) ρ density (kg.m-3) σ2 second moment variance

μ1 first moment

μ2 second moment

Acronyms ADM axial dispersion model CD catalytic distillation CFBR countercurrent fixed bed reactor RTD residence time distribution TIS tanks in series

Chapter 5: Bale Reactor Characterisation 250 5.8 Literature Cited Akbarnejad, M. M., A. A. Safekordi and S. Zarrinpashne “A study on the capacity of reactive distillation bale packings: experimental measurement, evaluation of existing models and preparation of a new model.” Ind. Eng. Chem. Res. 39: 3051-3058. (2000). Alicilar, A., M. Komurcu and A. Murathan “Air oxidation of aqueous cyanides in a countercurrent fixed bed reactor.” Korean J. Chem. Eng. 19(2): 273-276. (2002). Bennett, A. and F. Goodridge “Hydrodynamic and Mass Transfer studies in Packed Absorption Columns.” Trans. Instn. Chem. Engrs. 48: T232-T240. (1970). Bravo, J. L., J. A. Rocha and J. R. Fair “Pressure drop in structured packing.” Hydrocarbon processing 65(3): 45. (1986). Chow, C. W. K., D. E. Davey, D. E. Mulcahy and T. C. W. Yeow “Signal enhancement of potentiometric stripping analysis using digital signal processing.” Analytica Chimica Acta 307: 15-26. (1995). Edwards, M. F. and J. F. Richardson “Gas Dispersion in Packed beds.” Chem. Eng. Sci. 23: 109. (1968). Ellenberger, J. and R. Krishna “Countercurrent operation of structured catalytically packed distillation columns: pressure drop, holdup and mixing.” Chem. Eng. Sci. 54: 1339. (1999). Fogler, H. S. "Elements of Chemical reaction Engineering". New Jersey, Prentice Hall. (1999). Furzer, I. A. and R. W. Michell “Mixing in Trickle Flow through Packed Beds.” The Chem. Eng. J. 4: 53-63. (1972). Kister, H. Z. "Distillation Design". NY, McGraw-Hill Inc. (1992). Kunz, U. and U. Hoffmann (1995). Preparation of catalytic polymer/ceramic ionexchange packings for reactive distillation columns. Preparation of Catalyst VI. G. Poncelet, Elsevier Science B.V. Levenspiel, O. "Chemical Reaction Engineering". Brisbane, John Wiley & Sons. (1972). Marcias-Salinas, R. and J. R. Fair “Axial Mixing in Modern Packings, Gas and Liquid Phases: I. Single-Phase Flow.” AIChE Journal 45(2): 222-239. (1999). Marcias-Salinas, R. and J. R. Fair “Axial Mixing in Modern Packings, Gas and Liquid Phases: II. Two Phase Flow.” AIChE Journal 46(1): 79-91. (2000).

Chapter 5: Bale Reactor Characterisation 251 Mecklenburgh, J. C. and S. Hartland "The Theory of Backmixing". Sydney, John Wiley & Sons. (1975). Moritz, P. and H. Hasse “Fluid dynamics in Reactive Distillation packing Katapak-S.” Chemical Engineeering Science 54: 1367-1374. (1999). Otake, T. and E. Kunugita Chem. Eng. Tokyo 22: 144. (1958). Ramachandran, P. A. and R. V. Chaudhari "Three Phase Catalytic Reactors". New York, Gordon and Breach Science Publishers. (1983). Sater, V. E. and O. Levenspiel “Two-Phase Flow in Packed Beds.” I & EC Fundamentals 5(1): 86-92. (1966). Savitzky, A. and M. J. E. Golay “Smoothing and Differentiation of Data by Simplified Least Squares Procedures.” Analytical Chemistry 36(8): 1627-1639. (1964). Smith, L. A. J. "System for Separation iC4 from C4 Streams", 1980, US, 4215011 Stichlmair, J., J. L. Bravo and J. R. Fair “General model for the prediction of pressure drop and capacity of countercurrent gas/liquid packed columns.” Gas Separation Purification 3: 19. (1989). Stockar, U. v. and P. F. Cevey “Influence of the Physical Properties of the Liquid on Axial Dispersion in Packed Columns.” Ind. Eng. Chem. Process Des. Dev. 23: 717-724. (1984). Subawalla, H., J. C. Gonzalez, A. F. Seibert and F. J. R. “Capacity and Efficiency of Reactive Distillation Bale Packing: Modeling and Experimental Validation.” Ind. Eng. Chem. Res. 36: 3821-3832. (1997). Taylor, R. and R. Krishna “Modelling Reactive Distillation.” Chemical Engineeering Science 55: 5183-5229. (2000). Xu, X., Z. Zhihai and S. Tian “Study on Catalytic Distillation Processes. Part III: Prediction of Pressure Drop and Holdup in Catalyst Bed.” Trans. I Chem. E. 75: 625. (1997). Zhang, C. M., A. A. Adesina and M. S. Wainwright “Isobutene hydration in a countercurrent flow fixed bed reactor.” Chem. Eng. and Processing 43: 533-539. (2003).

Chapter 5: Bale Reactor Characterisation 252 6. Extractive Distillation over Bale Packing

This study was undertaken in order to evaluate the effectiveness of the reaction zone Bale packing in terms of TBA and water separation. As with most catalytic packings, bale packing separation efficiency tends to be low and it was considered desirable to determine how well the bale packing could contribute to the overall column. The column itself functioned as an extractive distillation and this mode of operation itself required validation. The loading limit and range of solvent flowrate also required confirmation.

An absorption representation of the extractive distillation process is adopted in this study. Thus, the solvent is fed as the liquid stream at the top of the column and the azeotrope as the gaseous feed stream at the bottom of the column. In this case water is preferentially absorbed from the gaseous feed. The bottom stream is taken off the column as opposed to being returned to the reboiler. The top vapour stream is condensed, but not returned to the column thus, the column was operated under the conditions of zero reflux. Strong parallels have been drawn between extractive distillation and the absorption process by various authors studying this type of homogeneous azeotropic distillation. Resa and Ruiz (1995) used the absorption representation to design equipment for evaluation of extractive distillation systems. Stichlamir and Herguijuela (1992) defined the first column of an extractive distillation train as an absorption column given that the low boiling azeotropic fraction is fed near the bottom and the high boiling entrainer is fed pure at the top. Additionally, both flowrates are higher than that of the reflux or reboil. The high boiling or heavy entrainer is of much lower relative volatility than the components being separated and thus does not appear in the top stream. Its prevalence in the bottom stream gives rise to the term extractive distillation, although at no point in the column are two actual liquid phases formed. Laroche et al (1992) presented evidence that in the case of homogeneous azeotropic distillation with a heavy entrainer, such as the separation of the water and ethanol azeotrope with ethylene glycol, the location of the entrainer feed relative to the azeotropic feed has a significant impact on separability. Separation becomes infeasible

Chapter 6: Extractive Distillation over Bale Packing 253 if the two feeds are introduced on the same tray because this eliminates the extractive section of the column. This documented result serves to strengthen the parallel between the middle extractive section of the column and absorption. The absorption representation allows direct focus upon the middle section of a standard extractive distillation column, that normally between the two feeds points, which most strongly influences the overall separation.

The aim of this work is to study the influence of solvent to azeotropic feed ratio and azeotropic feed molar flowrate on the performance of separation of TBA and water by extractive distillation with ethylene glycol over bale packing. The experimental results are analysed using phase-equilibrium information available in literature, Liu et al (1993), Yang and Wang (2002), and Smith, Van Ness and Abbott (1996). A further objective of this study is to validate the absorption type rig arrangement and operation at zero reflux for the purpose of investigation of catalytic extractive distillation development.

This chapter first presents the governing VLE, followed by a discussion of the experimental procedure and results. Furthermore, calculations are performed aimed at developing the proposed catalytic extractive distillation (CED) column.

Chapter 6: Extractive Distillation over Bale Packing 254 6.1 Governing VLE A literature review considering the availablity of parameters of suitable thermodynamic models for the prediction of activity coefficients of the significant study components water, isobutylene, TBA, ethylene glycol and MET, was conducted. Models considered suitable included the following: NRTL, Wilson, UNIQUAC and UNIFAC. Parameters of the NRTL and UNIQUAC models could not be found for any subset of the components of interest. Parameters for water, TBA and ethylene glycol could be found for the Wilson model, Liu et al. (1993) and Yang and Wang (2002). The UNIFAC model is a chemical group contribution model and given the relatively simple structures of the components all could be represented with parameter values available in Smith, Van Ness and Abbott (1996).

6.1-1 Wilson Model and Parameters Wilson (1964) proposed an expression for molar excess free energy using a logarithmic relation, which became the basis of the Wilson model for activity coefficients. The advantages of the model include its applicability for polar and non-polar solutions, the coefficients are temperature independent and as demonstrated by Oyer and Prausnitz (1965) mutlicomponent solutions require only parameters obtained from binary mixtures. The Wilson equation for mutlicomponent systems is given by:

NN ⎛⎞ xkkiΛ lnγ ijij=− ln ⎜⎟∑∑x Λ +1 − N (6.01) ⎝⎠jk==11 ∑ x jkjΛ j=1 where

Vji⎛⎞−λ j Λ=ij exp⎜⎟ (6.02) VRTi ⎝⎠

are Wilson parameters and λij is the binary interaction coefficient, which is obtained by fitting experimental data. For the non-symmetric binary interaction parameter Λij, generally Λij ≠ Λji, and further Λii = 1.

Chapter 6: Extractive Distillation over Bale Packing 255 The coefficients obtained by Liu et al (1993) for water, TBA and ethylene glycol were reported in Table 6.01. Expressions or values for molar volume, Vi, were not reported and yet are a significant part of the model.

Table 6.01 Wilson coefficients for TBA, Water and EG, Liu et al. (1993) TBA W EG TBA 7402.617 -379.009 W 5419.225 4143.67 EG 3627.66 -7460.31

The coefficients obtained by Yang and Wang (2002) for water and TBA were reported in Table 6.02. These values provide a means of comparison and validation of the values of Liu et al. (1993).

Table 6.02 Wilson coefficients for TBA and Water, Yang and Wang (2002) TBA W TBA 5916.8 W 5256.5

The use of the Wilson model requires knowledge of molar volume as a function of temperature. Molar volumes were calculated based upon density data available in Perry and regressed over the temperature range 290 to 423 K. A simple linear fit:

VaaT= 01+ (6.03)

was found adequate and the resulting coefficients were reported in Table 6.03.

Table 6.03 Molar Volume of TBA, Water and EG coefficients of Equation (6.03) TBA W EG a1 0.000132 1.01E-05 3.55E-05 a0 0.055463 0.014994 0.045172

The calculations were made using the spread sheet of Appendix A6.

Chapter 6: Extractive Distillation over Bale Packing 256 6.1-2 UNIFAC Model and Parameters The UNIFAC method relies on main and sub-group parameters, which are largely available for the components of interest. It requires that: 1. Pressure not greater than 5 bar be used. 2. Temperatures below 150 oC 3. No non-condensable components or electrolytes 4. Components must not contain more than ten functional groups The components and system under current consideration do not violate any of these recommendations. In the UNIFAC method the activity coefficient is expressed as combinatorial and residual parts:

CR lnγ ii=+ lnγγ ln i (6.04) where, following the notation of Smith, Van Ness and Abbott:

C ⎛⎞JJii lnγ iiii=− 1JJq + ln − 5⎜⎟ 1 − + ln (6.05) ⎝⎠LLii and

⎛⎞⎛⎞ββ lnγθR =−qe 1 ik −ln ik (6.06) ii⎜⎟∑⎜⎟ k ki ⎝⎠k ⎝⎠sskk

The quantities Ji and Li where defined:

ri Ji = (6.07) ∑ rxjj j

qi Li = (6.08) ∑ qxjj j

where r and q are obtained from experimentally derived subgroup parameters Rk and Qk by:

()i rRikk= ∑ν (6.09) k

Chapter 6: Extractive Distillation over Bale Packing 257

()i qQikk= ∑ν (6.10) k

In addition the following definitions apply:

()i ν kkQ eki = (6.11) qi

βik= ∑e miτ mk (6.12) m

∑ xiikiqe i θk = (6.13) ∑ x jjq j

skmmk= ∑θ τ (6.14) m

where τmk is calculated based upon main group interaction parameters amk by:

⎛⎞−amk τ mk = exp⎜⎟ (6.15) ⎝⎠T

The necessary subgroup parameters Rk and Qk as used were reported in Table 6.04 and the main group interaction parameters amk were reported in Table 6.05. The calculations were made using the spreadsheet presented in Appendix A5.

Table 6.04 UNIFAC Necessary Sub-Group Parameters

C CH CH2 CH3 OH H2O Q 0 0.228 0.54 0.848 1.2 1.4 R 0.2195 0.4469 0.6744 0.9011 1 0.92

Chapter 6: Extractive Distillation over Bale Packing 258 Table 6.05 UNIFAC Necessary Main Group Interaction Parameters amk C CH CH2 CH3 OH H2O C 0 0 0 0 986.5 1318 CH 0 0 0 0 986.5 1318

CH2 0 0 0 0 986.5 1318 CH3 0 0 0 0 986.5 1318 OH 156.4 156.4 156.4 156.4 0 353.5

H2O 300 300 300 300 -229.1 0

6.1-3 VLE Predicted For the calculation of the VLE governing the extractive distillation of water and TBA with the entrainer ethylene glycol the assumption was made that the vapour phase behaves as an ideal gas (setting ratio of fugacity coefficients, Φ, to unity). This allowed for modified Raoult’s law to be used, defined:

sat yiiiiPxP= γ (6.16)

VLE curves were calculated based upon the bubble point temperature approach detailed in Smith, Van Ness and Abbott (1996). The variable P, {xi} are used as setpoints and sat knowledge of Pi and γi behaviour in temperature and composition is required. The following equation lies at the heart of the approach: P Psat = (6.17) j ⎛⎞sat ⎛⎞xPiiγ i ∑⎜⎟⎜⎟sat i ⎝⎠Φij⎝⎠P

where Φi was set to unity. Values of yi and T are calculated iterating for convergence in T. The spreadsheet based calculator is given in Appendix A5 (UNIFAC) and A6 (Wilson).

Parameters for models describing the vapour pressure behaviour were generally plentiful. Those of the Antoine equation in truncated form:

sat ⎛⎞A2 PAi =−exp⎜⎟ 1 (6.18) ⎝⎠TA+ 3

Chapter 6: Extractive Distillation over Bale Packing 259 for water, isobutylene, TBA and ethylene glycol were sourced from Sinnott (1997) and reported in Table 6.06.

Table 6.06 Antoine Equation (6.18) Coefficients for TBA, Water EG and IB Comp. A1 A2 A3 TBA 16.85 2658.29 -95.50 H2O 18.30 3816.44 -46.13 EG 20.25 6022.18 -28.25 IB 16.95 2857.75 11.94

The binary x-y diagram for TBA/water in Figure 6.01 was generated using the models described. The agreement between experimental data and model prediction is good except at the dilute TBA end, where the UNIFAC model predicts a very high steep rise. This may be in response to the strong association of water molecules, which may be over predicted by the model. The azeotropes predicted at P = 1 atm were within moderate agreement: Wilson Liu xTBA.Az = 0.6754 and 353.7 K; Wilson Yang and Wang xTBA.Az = 0.6962 and 354.1 K; and UNIFAC xTBA.Az = 0.6107 and 353.4 K. These values compare well with data reported in the vapour liquid equilibrium data collection of

Gmehling and Onken (1977) of xTBA.Az = 0.6616 and 351.6 K. Thus the models are able to predict the most significant non-ideality of the TBA/water VLE scape. The Herington (1951) thermodynamic test for self-consistency of x-y data was applied. The model- generated VLE were found to be thermodynamically consistent. The calculated indicators (Wilson Liu et al 7.69, UNIFAC 3.29) were found to be within 10 units of the test criterion calculated to be 8.49. While the test is usually applied to experimental data, its use here did bolster confidence in the use of the models.

VLE obtained are presented in the x-y diagram form on a solvent free basis as commonly used for analysis of extractive distillation columns, Lynn and Hanson (1986). In applying such pseudo-binary VLE the assumption is made that the solvent mole fraction, xEG, remains nearly constant along the length of the column. The relative volatilities of both TBA and water are very large when compared to ethylene glycol. Due to its low volatility the solvent occurs almost exclusively in the liquid phase of the extractive distillation column.

Chapter 6: Extractive Distillation over Bale Packing 260 The solvent to feed ratio, Fr, is defined:

F Fr = SOLVENT (6.19) FAZEOTROPE

Additionally, the overall solvent mole fraction, xr, which is physically only valid in the range 0 < xr < 1, can be defined:

F xr = SOLVENT (6.20) FFSOLVENT+ AZEOTROPE

Combining and rearranging Equations (6.19) and (6.20) leads to:

xr Fr = (6.21) 1− xr

Thus, Fr → ∞ as xr → 1. The effect of xr was explored using the Wilson and UNIFAC models, Figure 6.02 (a) and (b). The Wilson model with parameters of Liu et al (1993) predicts that the selectivity of ethylene glycol is very strong with the x-y diagram almost completely opened up at xr = 0.33 (Fr = 0.5). The UNIFAC model predicts much more pinching behaviour, which is slowly overcome with the addition of ethylene glycol, however little opening up and broadening of the distillation space defined by the equilibrium and 45o line. The broader this space the fewer equilibrium stages are required to achieve the same degree of separation of TBA and water. The Wilson model predicts that at the water end of the x-y diagram selectivity is marginally improved and the curve opens up. Conversely, the UNIFAC model suggests that solvent addition is detrimental to separation at this stripping end. The curve becomes symmetric by closing in at the water end. It is possible that the UNIFAC model struggles with water rich mixtures.

Chapter 6: Extractive Distillation over Bale Packing 261

1

0.9

0.8

0.7

0.6

TBA 0.5 y 0.4

0.3

0.2 Liu et al. (1993) Yang & Wang (2002) 0.1 UNIFAC 0 0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

xTBA

Figure 6.01 Comparison of binary x-y data for the TBA water system generated using the Wilson and UNIFAC models showing the presence of an azeotrope

occurring at approximately xTBA 0.60 to 0.70.

Chapter 6: Extractive Distillation over Bale Packing 262 1 xr 0 0.8 xr 0.33 xr 0.5 0.6 xr 0.67 TBA

y' xr 0.75 0.4 xr 0.8

0.2

0 0 0.2 0.4 0.6 0.8 1 (a)

x'TBA

1

xr 0 0.8 xr 0.33 xr 0.50 0.6 xr 0.67

TBA xr 0.75 y' 0.4 xr 0.80

0.2

0 (b) 0 0.2 0.4 0.6 0.8 1

x'TBA

Figure 6.02 X’-Y’, VLE data presented on a solvent free basis demonstrating the ability of ethylene glycol to cross the TBA/water azeotrope. Generated using (a) Wilson model with parameters of Liu et al. and (b) UNIFAC model.

Chapter 6: Extractive Distillation over Bale Packing 263 6.1-4 Relative Volatilities and Selectivity Selectivity is defined as the ratio of the relative volatilities of the key components of the mixture, which are to be separated, in the presence of the solvent to their relative volatilities before the addition of the solvent. Thus, for key components i, j and solvent s, selectivity Sij is defined, Momoh (1991):

αijs Sij = (6.22) αij where the relative volatilities are defined by:

o γ iif θ j αij = o (6.23) γ ijif θ

o At low to moderate pressures and temperatures the standard state fugacity, fi , can be sat approximated by pure component vapour pressure, Pi , and the rate of the vapour phase fugacity coefficient, θ, is usually close to 1.0. Thus:

sat γ iiP αij = sat (6.24) γ jjP

The relative volatilities of TBA and water in the absence and presence of the solvent ethylene glycol were calculated according to the Wilson and UNIFAC models and are presented in Figure 6.03 (a) and (b). Both the models predict that ethylene glycol improves the relative volatility for mid to high x’TBA to well above unity thus clearing the azeotrope and azeotropic behaviour. In each case additional solvent leads to an increase in relative volatility within this region of 0.4 < x’TBA < 0.9. and The Wilson model, with parameters gained from experiments conducted directly on the system, gives relative volatilities which suggest that the solvent is 4 to 5 times more effective than that predicted by the UNIFAC model. In the lower TBA range the UNIFAC model predicts a decrease in relative volatility.

Chapter 6: Extractive Distillation over Bale Packing 264 20 19 18 xr 0 xr 0.1 xr 0.3 17 xr 0.5 xr 0.7 xr 0.9 16 15 14 13 12 11 10 TBA.W 9 α 8 7 6 5 4 3 2 1 0 0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

x'TBA (a)

13 12 xr 0 xr 0.1 xr 0.3 11 xr 0.5 xr 0.7 xr 0.9 10 9 8 7

TBA.W 6 α 5 4 3 2 1 0 0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

x'TBA (b)

Figure 6.03 Relative volatilities predicted by (a) Wilson and (b) UNIFAC models described.

Chapter 6: Extractive Distillation over Bale Packing 265 Selectivity as a function of xr using the VLE calculations outlined was determined at the azeotrope composition predicted. The relative volatility at the azeotrope is 1.0, thus the selectivity and relative volatility are equal. The predicted selectivities, Figure 6.04, are consistent with the qualitative observations made for the x-y diagrams. These selectivities diverged with the UNIFAC model giving low values, about 2.4. The Wilson model gave values, which increased exponentially towards 11.96 for near pure solvent. Given the range of values it is possible that the UNIFAC model under predicts and the experimentally derived values over predict the true selectivity. The difference in predicted selectivities were noted with care given the sensitivity of homogenous azeotropic distillation to thermodynamic data inaccuracy and the inherent design uncertainty noted by Laroche et al (1992).

12 Wilson 10 UNIFAC

8

6 TBA.W S 4

2

0 0 0.2 0.4 0.6 0.8

xEG

Figure 6.04 Selectivity calculated by the Wilson Model at x’TBA = 0.6754 and

UNIFAC at x’TBA = 0.6107, the respective predicted solvent free azeotropes.

Chapter 6: Extractive Distillation over Bale Packing 266 6.1-5 Minimum Solvent Required The gradual improvement in relative volatility and the eventual clearance of unity suggest that there is a minimum amount of solvent, which is required to overcome the azeotrope. This minimum required ratio of solvent can be estimated using the expression suggested by Pretel et al (1994):

xrmin ,1.0αij = (6.25) xS .min

At xA or xB = 0 depending on the type of extractive distillation, that of promoting or reversing natural tendency, xr is iterated until αBA = 1. This expression implies that the azeotrope will shift to the pure TBA end and almost off the binary curve.

This method can be actualised directly by VLE calculation such that the correct T for a given P (1 atm) is used and the ratio of vapour pressures is correct. The value of αij was calculated at x’TBA = 0.9999 (dilute x’W). A search was conducted using a method of -04 bisection in order to determine xr.min for a tolerance of 10 . The results are shown in Table 6.07 for the Wilson and UNIFAC models. It was determined that the minimum solvent required, xr.min by the UNIFAC model was xrmin = 0.4761 and by the Wilson model xrmin = 0.0625. The estimates are very different and as found from a comparison of VLE curves the parameters of Liu et al (1993) suggest that ethylene glycol is highly effective.

Table 6.07 Determination of xrmin using Wilson and UNIFAC Models for the Extractive Distillation of Water and TBA.

Method Bisection xr αij Wilson 1 0.0100 0.6894 2 0.3300 4.0450 3 0.1700 1.9133 4 0.0900 1.1967 5 0.0500 0.9188 6 0.0700 1.0514 7 0.0600 0.9836 8 0.0650 1.0171

xrmin : 9 0.0625 1.0002

UNIFAC 1 0.3300 0.8096 2 0.5000 1.0354

Chapter 6: Extractive Distillation over Bale Packing 267 3 0.4150 0.9151 4 0.4575 0.9732 5 0.4788 1.0038 6 0.4681 0.9882 7 0.4734 0.9954

xrmin : 8 0.4761 0.9999

While this approach has a sound basis it was found that care should be taken to ensure that the solution is not a pinch or pseudo-azeotrope. Generation of VLE allows for determination of the shape of the curve. In particular, whether it is convex or concave in nature, when the azeotrope is broken. The difference between y’TBA predicted by VLE o calculation and that given by the 45 line of y’TBA = x’TBA was defined as β and is plotted as a function of x’TBA for the azeotropic end of the x-t plot in Figure 6.05 (a) and (b).

It can be seen that the Wilson estimated value is marginally smaller than that which actually clears the azeotrope at approximately xr = 0.072, obtained by trial and error. Moreover, the curve is concave creating a pinch with the 45o line. For the Wilson estimate the pinch is well diminished by xr = 0.165 and the curve becomes convex. Similarly, the UNIFAC estimate of xr = 0.48 creates a pinch which is overcome by xr =

0.75 (Fr = 3). A second order polynomial fit of the data shows that a1 and a2, for xr = 0.75, both become negative and the differential increases in absolute magnitude showing that the point (0,1) is approached from above. Thus xrmin.non-pinch = 0.165 and xrmin.non-pinch = 0.75 as predicted by the Wilson and UNIFAC models respectively. These values are twice those predicted as being the minium required to break the azeotrope.

The above mentioned calculations demonstrate that the minimum amount of solvent required to break the azeotrope can lead to a VLE curve with a pinch. Such a pinch will see little improvement in purity with additional stages. Thus, greater care is required when predicting the minimum amount of solvent. An understanding of the behaviour of the VLE will assist in the analysis of extractive distillation experiments conducted over bale packing and described in subsequent sections. Knowledge of bounds, such as minimum solvent fraction, minium non-pinching solvent fraction and maximum potential, help reference and define the experimental set points.

Chapter 6: Extractive Distillation over Bale Packing 268

0.2 xr 0 0.15 xr 0.063 xr 0.165 xr 0.33 0.1 β

0.05

0

-0.05 (a) 0.60.811.2

x'TBA

0.14 xr 0.33 0.12 xr 0.48 0.1 xr 0.50 xr 0.67 0.08 xr 0.75

β 0.06 xr 0.80

0.04

0.02

0

-0.02 0.6 0.7 0.8 0.9 1 1.1 (b)

x'TBA

Figure 6.05 The approach of the equilibrium line to the 45o line for xr, (a) Wilson and (b) UNIFAC models

Chapter 6: Extractive Distillation over Bale Packing 269 6.2 Experimental In the following section the equipment, experimental setup, procedure, and experimental design are detailed, for the extractive distillation of the TBA and water with ethylene glycol as solvent conducted using inert Bale packing.

6.2-1 Distillation Equipment The bale packed CFBR reactive zone characterised in Chapter 5, was incorporated into a distillation column specifically designed for the development of extractive distillation. Specialist glass was designed and fabricated allowing for the extractive distillation to be operated essentially as an adsorption column. This was achieved according to the design of reboiler bypass and semi-batch operation proposed by Resa et al (1995). This design has previously been used to dynamically determine the isobaric VLE of acetone/propyl ether and isopropyl ether/propyl ether systems, as well as well as the effectiveness of extractive distillation of these systems with a wide range of entrainers, Resa et al (2000). The bottoms are not allowed to return to the reboiler by use of a vapour liquid separator, Figure 6.06 glassware piece 5. Vapour is allowed to rise up the central snorkel, while liquid is taken off. The reboiler can be refilled with the lower extended connection or the liquid taken off above the split can be returned to the column. The small volume created by the height of the snorkel was calibrated with a scale mounted on the outside. This was used to measure column throughput and bottom flowrates. All glassware was fabricated from Pyrex and individual pieces were joined using QuickFit ground glass connections, Figure 6.06. Pieces 1 and 2 were fabricated for use with the complete CED design and for extractive distillation experiments the condenser depicted (coil condenser, QF 55/44) in Figure 6.06 was used and mounted directly upon piece 3. It was also equipped with a vapour liquid separator. The internal diameter of the column was 45 mm. The column ports were 3/8” glass tube and Swagelok connectors with teflon ferrules were used to connect tubing, septums and thermocouple fittings to these.

Chapter 6: Extractive Distillation over Bale Packing 270 id 45 1. 840 50/42

1. 2.

600 teflon

50/42 3. 240 2.

120 4. 90

55/44

4. 300 55/44

225 3.

115 5. 50/42

50/42 50/42 5. 240

12 6. 125 75

55/44

Figure 6.06 QuickFit Distillation Glassware: 1. Condenser Connection; 2. Reflux Section; 3. Solvent Feed Section; 4. Jacketed Bale packed Section; 5. Reboiler bypass; and 6. Reboiler

Chapter 6: Extractive Distillation over Bale Packing 271 6.2-2 Analytical Procedure The determination of water in TBA, and TBA and water in ethylene glycol were carried out using gas chromatography (GC) and liquid phase manual injection. The components of the mixtures analysed were of comparable concentrations. A Shimadzu GC-8A equipped with TCD detector and Shimadzu C-R8A integrator was employed. It was fitted with a packed stainless steel column recommended and manufactured by Alltech, with packing HaySep Q of properties: particle size 60/80 mesh, column diameter 1/8”, length 8’ and Tmax 523 K. Helium was employed as the carrier gas and the liquid injection volume was 2 μL delivered by micro-syringe. All other GC and Integrator settings are reported in Table 6.08.

Table 6.08 GC and Integrator settings for analysis of water and TBA within water/TBA/Ethylene Glycol mixtures. GC Integrator Variable Value Variable Value Carrier 1, Helium 220 kPa Attenuation 7 Carrier 2, Helium 50 kPa Stop Time 10 min Inj/Det Temp. 240 oC Peak Width 1 Col. Temp. 210 oC Min. Area 50 Current 80 mA Polarity Plus Attenuation 1

Extractive distillation column samples taken were immediately sealed and rested allowing them to cool down to room temperature. These samples were sub-sampled and combined with a given volume of the internal standard ethanol (60 μL sample to 40 μL

EtOH). The calibration of concentration versus area (Ai) obtained are reported in Table 6.09.

Table 6.09 Calibration of area versus concentration for analysis of TBA and water. Analysis Calibration R2 Eq. No.

W in ⎛⎞A 0.9903 (6.26) EG/TBA W CW = 22.4658⎜⎟

⎝⎠AEtOH TBA in ⎛⎞ 0.9773 (6.27) EG/Water ATBA CTBA = 9.1008⎜⎟

⎝⎠AEtOH

Chapter 6: Extractive Distillation over Bale Packing 272 6.2-3 Experimental Procedure The extractive distillation rig is depicted in Figure 6.07. The experiments were performed in a column employing bale packing filled with inert glass spheres as used for the fluid dynamic characterisation. The number of Bales was kept at five as this would be the number used for CED runs and was the number of bales that could be held by the reactive zone.

For extractive distillation the reboiler was charged with azeotrope prepared in the same column at atmospheric conditions using a sub-azeotropic feed. The column was started up under conditions of total reflux. Once the azeotrope had boiled, appeared at the top of the column, and condensed liquid had returned to the bottom, the ethylene glycol could be fed and the experiment commenced for a given boilup rate and solvent feed rate. The experiments were performed under conditions of zero reflux, giving a stream arrangement akin to absorption. The concentrations of TBA and water were determined by GC analysis using TCD and manually injected liquid phase samples as detailed in Section 6.2-2.

Generally a solvent feed temperature in the range of the boiling points of the components to be separated is chosen for extractive distillation. As such, a solvent feed temperature between 353 and 373 K, the atmospheric boiling points of TBA and water respectively was required. In this case 373 K was selected in order to maximise top vapour flowrate and allow for ease of sampling and flow measurement. It was considered undesirable for the solvent to merely condense the azeotrope, instead preferential absorption of water was sought. Furthermore, in the CED final design higher temperatures would be favoured, giving according to the relative reaction rate findings of Chapter 4, higher selectivity in TBA and faster reaction rates. Solvent feed temperature is a significant but often overlooked variable in extractive distillation, Momoh (1991). As with most variables in extractive distillation, solvent feed temperature leads to a trade off between recovery and product purity.

For extractive distillation reflux ratio is often found to be an insensitive variable, fixed at some low level, Ruiz et al (1997). There are several potential advantages to operating at zero reflux or minimal liquid return offered by the column/condenser connection

Chapter 6: Extractive Distillation over Bale Packing 273 zone. Firstly, a reduction in hardware associated with flow splitting and measurement as well as reflux preheating. Generally the distillate stream is subcooled and requires preheating before being fed back to the column. The task of reheating a small, often discontinuous, flow is difficult and fluctuations can lead to strong dynamic disturbances, Kumar et al (1984). Reflux, also serves to dilute the solvent thus decreasing its selectivity. Most significantly, if extractive distillation is combined with catalytic distillation, conditions of zero reflux mean less product recycle to the reaction zone. It is anticipated that the use of solvent will reduce the need for product recycle. If it is possible to work with low reflux then smaller amounts of product will be returned to the reaction zone and the equilibrium shifting effect will be stronger. Generally, gaining higher purity with reflux comes at the expense of recovery. This is especially true if reflux is making up for an inadequate number of stages.

A disadvantage of operation under zero reflux is that standard distillation analysis procedure developed for total or partial reflux cannot be utilised. Such a standard closed system method has previously been used with bale packing by Subawalla, Gonzalez, Seibert and Fair (1997), for the chemical systems of acetone/MEK and cyclohexane/n- heptane. However, extractive distillation fails at infinite reflux (the column does not perform any separation) and often has a maximum efficiency at finite reflux, Laroche et al (1992). Thus, standard methods of analysis give at best an approximate indication of efficiency.

The temperature distribution of the column was monitored by an even placement of minerally insulated K-type thermocouples. The column was insulated with three layers of specialist fibreglass tape as used for steam supplies. The rate of boilup was controlled directly by a heating mantle power regulator. The rate of solvent feed was controlled by a variable speed peristaltic pump and monitored by a rotameter. The column was maintained at atmospheric pressure by adjusting the condenser cooling water such that the level of an indicating manometer system was balanced.

Chapter 6: Extractive Distillation over Bale Packing 274 [01] [02] [03] T

1 [04] [05]

6 2 7 [06]

T 3

[07] [08] [09] T 8 4 ETHYLENE GLYCOL TANK [10] [11] 5 SILICON OIL TANK

[13]

[12] [14]

BOTTOMS TANK

Figure 6.07 Extractive Distillation Rig

Chapter 6: Extractive Distillation over Bale Packing 275 Key to Figure 6.07 Extractive Distillation Rig also used in Distillation Liquid Dynamics [01] Temperature Indicator and Thermocouple Selector Switch, Shinko Controller, RS Selector. [02] Condenser: Large Coil, QuickFit B55. [03] Pressure Relief/Indicator. [04] Column Reflux Assembly. [05] Cooling water rotameter. [06] Entrainer Feed Tank: Braun Bath Heater, SS bath. [07] Entrainer Feed Rotameter: Planton, Duff and Machintosh, Model 48, universal, with scale marked for 2 to 25 L/min Air. [08] Column: Jacketed 45 mm id, 60 mm od. [09] Jacket Heating Oil Bath: Bath Heater and Recirculator, Haake DC 5; Insulated SS tank. [10] Liquid Feed Pump: Masterflex Peristaltic Pump, Variable Speed Variable Speed Modular Drive, Model 7016 pump head, L/S 16 Viton Tubing, Cole Parmer. [11] Column Reboiler Bypass QuickFit Fitting: also used as calibrated volume to measure bottom liquid flow [12] Liquid Waste Tank: 20 L polypropylene waste container. [13] Glass Reboiler, 10 L or 5 L, QuickFit single opening. [14] Heating Mantle, 10 L, with Power regulator.

Extractive Distillation Experimental Procedure 1. Prepare GC-8A TCD for analysis. 2. The column [08] should be packed with inert bale packing containing glass particles. 3. Prepare EG Tank Charge, make sure there is sufficient room in the waste container, prepare a spare if necessary. 4. Charge the reboiler [13] with the TBA in water azeotrope feed, prepared earlier in same column. 5. Turn on the condenser water [05] 2 to 3 L/min. Failure to turn on the condenser water can result in column damage and condenser tubing damage.

Chapter 6: Extractive Distillation over Bale Packing 276 6. Check that the reflux is turned to the desired position [04] and that the pressure relief/indicator is in place [03]. 7. Preheat the EG to the desired temperature [06], typically 80 oC. 8. Turn on the reboiler [13] and set to the maximum 100% power. This setting should later be changed to the desired experimental value once the azeotrope boils. 9. Turn on the Silicon Oil Recirculator and set to the desired value, process + 15 oC. 10. Once the azeotrope has boiled and the reboiler has been adjusted wait till the azeotrope begins to condense within the condenser and begins to return to the column. 11. Turn on the EG pump [10] and set the flow rate to that required using the rotameter [07], turn the three way valve towards the column to begin the experimental run. The solvent flowrate will need to be adjusted at this point to compensate for the higher-pressure drop of pumping into the column. 12. Monitor and open the bottoms bypass [11]. Turn of the flow and measure the bottoms flow rate using the calibrated volume and a stopwatch. 13. Take the temperature profile at regular intervals using the selector switch and temperature indicator [01]. Watch this profile by plotting for the approach to steady state. 14. Take top and bottom samples every 10 minutes once the temperature stabilises take at least five samples looking for steady state. 15. To end run turn off the EG, top up the azeotrope and turn to max boil up rate and wash down the column. Set the new setpoints of the next run. 16. For complete shutdown turn off reboiler and preheaters. Allow the column to cool down and then turn off cooling water.

Chapter 6: Extractive Distillation over Bale Packing 277 6.2-4 Experimental Variables As a preliminary phase to the catalytic extractive distillation, this study was carried out with a solvent feed temperature of 373 K with zero reflux using an azeotropic feed composition in a 5 stage (bale) bed.

Azeotrope loading in the form of boilup rate and solvent feed rate were varied across their full practical ranges, limited by flooding and drying up of distillate due to reverse entrainment. Starting at a middle boilup rate of 0.57 mol.min-1 of azeotrope, obtained at 50 % of 2400 Watt, the boilup rate was varied upwards and downwards by 240-Watt increments. The solvent feed rate was varied to give solvent to azeotropic feed ratios of Fr = 0.5, 1, 2, 3, and 4. The lower ratio achieved was above both the minimum solvent required Frmin = 0.07 and minimum solvent to move away from pinching activity

Frmin.nonpinch = 0.20 as determined in Section 6.1-5 according to the Wilson model.

Chapter 6: Extractive Distillation over Bale Packing 278 6.3 Extractive Distillation Results The results are presented in Table 6.10. The streams referred to are identified in the stream diagram for extractive distillation experiments of Figure 6.08. The azeotrope was successfully broken with the column operating at zero reflux as demonstrated by the column of distillate TBA mole fraction data x’TBA.3, where solvent free values are well over the azeotrope composition of 0.65. The ethylene glycol mole fractions of xr = 0.4 to 0.75 were achieved for runs 7 to 18 of middle boilup range. However, for the lower boilup rate (0.49 mol.min-1) the distillate tended to run dry with high solvent rates and for the high boilup rate (0.82 mol.min-1) the solvent was entrained (appeared in distillate) limiting the range of experimentation.

Condenser 3. Distillate/Product

F3 (TBA,W)

2. Solvent F (EG) 2 Extractive Distillation Column

Bale Packing

1. Azeotrope 4. Bottoms

F1 (TBA,W) F4 (TBA,W,EG)

Figure 6.08 Extractive Distillation Stream Diagram

Chapter 6: Extractive Distillation over Bale Packing 279 Table 6.10 Extractive Distillation of TBA and Water with Ethylene Glycol Main Results Stream: 1 2 3 Feed Ratio Feed Frac. Boilup Solvent Distillate Purity Recovery

Fr xr F1 F2 Q3 CW.3 CTBA.3 PTBA RTBA Run ID mol.min-1 mol.min-1 mol.min-1 mL.min-1 mol.L-1 mol.L-1 1 0.84 0.46 0.49 0.41 14.88 1.66 10.50 0.86 0.32 2 1.08 0.52 0.49 0.53 11.39 1.64 10.94 0.87 0.26 3 1.64 0.62 0.49 0.80 1.78 0.72 10.51 0.94 0.04 4 1.96 0.66 0.49 0.95 - - - - - 5 2.68 0.73 0.49 1.30 - - - - - 6 0.72 0.42 0.57 0.41 33.04 1.27 10.21 0.89 0.59 7 1.15 0.54 0.57 0.66 23.41 0.86 10.14 0.92 0.42 8 1.67 0.63 0.57 0.95 30.16 0.80 9.99 0.93 0.53 9 1.97 0.66 0.57 1.12 6.81 0.69 10.92 0.94 0.13 10 2.63 0.72 0.57 1.50 1.84 0.90 10.65 0.92 0.03 11 0.58 0.37 0.71 0.41 33.64 2.30 10.01 0.81 0.47 12 0.92 0.48 0.71 0.66 32.00 2.51 10.24 0.80 0.46 13 1.58 0.61 0.71 1.12 25.67 1.79 10.56 0.86 0.38 14 2.11 0.68 0.71 1.50 16.86 0.89 11.01 0.93 0.26 15 2.71 0.73 0.71 1.92 9.22 0.75 10.86 0.94 0.14 16 0.64 0.39 0.82 0.53 33.24 2.40 10.15 0.81 - 17 0.97 0.49 0.82 0.80 39.27 2.49 10.67 0.81 - 18 1.58 0.61 0.82 1.30 24.65 1.90 10.09 0.84 -

Chapter 6: Extractive Distillation over Bale Packing 280 Table 6.11 Extractive Distillation of TBA and Water with Ethylene Glycol, Column Temperature Profiles Run ID 5 4 3 2 1 2 315 354 371 361 353 3 313 357 369 363 353 4 320 356 374 370 358

8 359 373 367 356 355 9 362 376 369 358 354 10 361 382 369 357 354 11 355 377 370 361 355 12 355 377 372 369 355

14 362 378 364 356 358 15 361 378 365 357 359 16 361 374 369 361 357 17 361 381 370 358 357 18 361 375 371 365 357

20 355 351 351 352 352 21 362 370 378 367 355 22 362 380 378 368 356

Chapter 6: Extractive Distillation over Bale Packing 281 6.3-1 Range of Operation

The experimental design approach generated ReG in the range 366 to 620 and ReL in the range 6 to 33. According to the loading Equation (5.41):

−0.78 ReG, Loading = 21580ReL (5.41) developed in Chapter 5 Section 5.06, all runs were conducted below the loading point in -1 the pre-loading regime. Runs 4 and 5, of azeotrope loading F1 = 0.49 mol.min and xr = 0.66 and 0.73 suffered from reverse entrainment and the product stream dried out. Runs -1 19 and 20, F1 = 0.82 mol.min and xr = 0.7 and 0.74 suffered from severe solvent entrainment.

6.3-2 Temperature Profiles Temperature was measured at the various points shown in Figure 6.07, using minerally insulated K-type thermocouples and reported in Table 6.11. The temperature profiles obtained did not vary greatly with azeotrope loading, but were a strong function of solvent flow rate and temperature. The typical variation in solvent feed rate is shown in Figure 6.09. The sampling points were approximately equidistant. The first sample point was the reboiler, while the last was the top of the column. The solvent feed was located at sample point 4.

The profile in Figure 6.09 clearly shows that the maximum temperature was always at the solvent feed location and dropped towards reboiler and the reflux stream temperature on either side of the peak irrespective of the ethylene glycol content. As may be expected, during startup at total reflux, when there was no solvent feed the temperature profile was essentially flat. This suggests that the heat input from the reboiler was probably equal to the heat removed by the condenser returning the reflux stream.

In general, for a given feed, the difference between the maximum and minimum temperature in the bale packed zone was 15 K. In the final CED application the feed temperature can be optimised and most likely a flatter temperature distribution obtained

Chapter 6: Extractive Distillation over Bale Packing 282 reducing the difference across the reaction zone and obtaining near isothermal conditions, which is atypical of catalytic distillation. The solvent also gives some scope to manipulate reaction temperature independently of column pressure, which is also less than typical for catalytic distillation.

5

xr 0.42 xr 0.54 4 xr 0.63 xr 0.66 xr 0.72

3

2 Sampling Point ( / 0.3 m)

1 320 340 360 380 400 T (K)

Figure 6.09 Extractive Distillation Temperature Profile at Az 0.57 mol.min-1.

Chapter 6: Extractive Distillation over Bale Packing 283 6.3-3 Separation by Extractive Distillation The feed azeotrope composition of TBA and water has an experimentally determined mean value of xTBA = 0.6634 ± 0.16. This azeotrope was successfully crossed using the solvent ethylene glycol under conditions of semi-batch extractive distillation (cf. Figure 6.10). The distillate or stream 3 TBA mole fraction, which can be referred to as product purity PTBA, approached 0.95. The purity did seem to plateau, as particularly evident of the data collected at F1.Az. Given operation over five bales and a short bed of Raschig rings this is likely to be experimental evidence of the pinching behaviour predicted by both the Wilson and UNIFAC models. Higher azeotrope loading, which corresponds to higher vapour velocities, did not have a direct influence on purity as may be expected for conventional distillation where efficiency increases with vapour phase loading.

The recovery of TBA was defined:

Fx3.3TBA RTBA = (6.28) Fx1.1TBA and the values generated were plotted in Figure 6.11. It was found that for operation at zero reflux in an absorption type setup, without a stripping section, the best recovery possible was RTBA = 0.6. The recovery dropped away sharply as solvent fraction was increased. It increased with boilup rate or azeotrope loading. This suggested heavy TBA entrainment at the bottom of the experimental setup, which is to be expected given the location of the feed. This, however, should not be a concern for catalytic extractive distillation runs where the TBA would be generated throughout the bale bed.

Chapter 6: Extractive Distillation over Bale Packing 284 1.00 0.95 0.90 0.85 TBA.3

: x 0.80

TBA 0.75 P F1 = 0.49 mol/min 0.70 F1 = 0.57 mol/min 0.65 F1 = 0.71 mol/min 0.60 0.30 0.40 0.50 0.60 0.70 0.80 xr

Figure 6.10 Purity gained by extractive distillation as a function of xEG for different azeotrope loading Az: 1. 0.49, 2. 0.57, 3. 0.71, and 4. 0.82 mol.min-1

1 0.9 F1 = 0.48 mol/min 0.8 F1 = 0.57 mol/min 0.7 F1 = 0.71 mol/min 0.6

TBA 0.5 R 0.4 0.3 0.2 0.1 0 0.30 0.40 0.50 0.60 0.70 0.80 xr

Figure 6.11 Recovery in extractive distillation as a function of xEG for different azeotrope loading Az: 1. 0.49, 2. 0.57, and 3. 0.71 mol.min-1

Chapter 6: Extractive Distillation over Bale Packing 285 A plot of recovery versus purity gives some indication of the optimum region of the current experimental design, Figure 6.12. The point closest to the intersection of complete purity and recovery (1,1), corresponds to an azeotrope load, F1 = 0.57 mol.min-1 and a solvent to feed ratio of Fr = 1.58 mol.mol-1 or xr = 0.42, and corresponds to optimum purity and recovery. The plot indicates that the zero reflux setup is biased towards purity and poor recovery itself may limit the attainable purity due to completely absorbed vapour phase.

1.0 0.9 0.8 0.7 0.6

TBA 0.5 P 0.4 0.3 F1 = 0.48 mol/min 0.2 F1 = 0.57 mol/min 0.1 F1 = 0.71 mol/min 0.0 0 0.2 0.4 0.6 0.8 1

RTBA

Figure 6.12 Purity against Recovery for the extractive distillation study conducted with different azeotrope loading Az: 1. 0.49, 2. 0.57, and 3. 0.71 mol.min-1

Chapter 6: Extractive Distillation over Bale Packing 286 6.3-4 Theoretical Plates Achieved The experimental results are analysed for the number of theoretical stages achieved with respect to the governing VLE detailed in Section 6.1. The experimentally obtained VLE of Liu et al (1993) correlated by the Wilson model were employed as opposed to the VLE predicted by the UNIFAC thermodynamic group contribution model. The high purities obtained and the responsiveness of the system suggested that the true relative volatilities are closer to those of the Wilson model.

The number of equilibrium stages attained over the 0.6 m of inert Bale packing was estimated using a modified version of the conventional graphical approach of McCabe and Thiele detailed in Fair (1987). The McCabe-Thiele method has been previously applied to extractive distillation using the pseudo-binary approximation, Atkins and Boyer (1949) and Meirelles and Telis (1992). These applications were carried out under conditions of finite reflux, where the assumption of constant vapour and liquid pseudo- binary molar flowrates along the length of each column section can be justified. Given this assumption the operating lines obtained were linear. The stream configuration of the current study, with zero reflux employed, resembled closely absorption. Under such conditions the entrainer initially free of water or TBA, absorbs increasingly more water and TBA along the length of the column, thus the ratio of pseudo-binary vapour and liquid flowrates changes considerably along the length of the column and the operating line for this arrangement is non-linear. While the entrainer absorbs both species its action upon the liquid phase activities is such that the relative volatility between TBA and water is increased and separation of water and TBA beyond the azeotrope is achieved as TBA is promoted into the vapour phase.

While, it was not possible to take samples along the length of the column given the structure of bale packing, the operating line could be approximated given knowledge of the inlet and outlet conditions. The slope of the rectifying operating line, mopp, is defined as: dy L m = = (6.29) opp dx V

Chapter 6: Extractive Distillation over Bale Packing 287 where L and V are the liquid and vapour molar flowrates. In terms of pseudo binary flowrates at the base of the column referring to Figure 6.08 the slope is given as:

F4′ mopp.1 = (6.30) F1′

where F1’ and F4’ are given on a solvent free basis. At the top of the column for all runs zero reflux was returned and solvent was free of water and TBA thus:

F2′ 0 mopp.2 = ==0 (6.31) FF33′ ′

The operating line for each run was estimated as a quadratic equation:

2 y = ax120++ ax a (6.32)

satisfying Equation (6.30) and (6.31) and passing through the distillate and bottoms compositions. The equilibrium line was calculated based upon xr of the particular run.

The relevant section of the VLE curve, typically in the range 0.3 ≤ x’TBA ≤ 1.0, was represented using a quadratic fit allowing for ease of calculation.

The x’-y’ equilibrium curves and operating curves obtained are provided in Figure 6.13 -1 -1 for F1 = 0.49 mol.min , Figure 6.14 for 0.57 mol.min , and Figure 6.15 for 0.71 mol.min-1. Each figure shows that as the molar fraction of solvent, xr, is increased two significant effects occur. Firstly, the distillate point slides along the 45o line towards (1,1) in correspondence with increased product purity. Secondly, the critical space between the operating lines and the equilibrium lines for each run becomes broader and the difficulty of separation is diminished. The lines tend to intersect at low xr, which is further evidence of the pinching effect demonstrated in Section 6.1-5 pertaining to the evaluation of minimum solvent required. With further increase in xr they become almost parallel (equispaced) indicating comfortable separation with a finite number of stages. Finally, with increased xr the operating and equilibrium lines become highly divergent. This divergence is created by a high degree of TBA reverse or downward

Chapter 6: Extractive Distillation over Bale Packing 288 o entrainment, which leads to the value of mopp.1 approaching unity and a 45 angle. This high reverse entrainment over the relatively short column obviously results in poor recovery. The gradual change from pinching lines, to parallel to divergent with increased xr suggests that there is a xr value at which the curves go from being near parallel to divergent. This value approximately occurs at xr = 0.52 for F1 = 0.49 -1 mol.min , xr = 0.54 for F1 = 0.57, and xr = 0.61 for F1 = 0.71. These values correspond to the commencement of severe reverse entrainment. Moreover, the proposition can be made that for a given length of test column the more divergent the operating and equilibrium curves the greater an improvement in separation, both purity and recovery, can be made with an increase in length. However, this also suggests that for a given xr there is likely to be a length at which the curves become parallel or convergent and very little improvement can be gained especially in purity of the product with increased number of stages. These observations require further investigation and may be a particular characteristic of the current test system.

The number of theoretical plates achieved was determined in the conventional manner by stepping of the stages away from the distillate composition, with each vertical line constructed counting as one theoretical stage achieved. The values obtained are presented in Figure 6.16. With increased xr the number of stages attained decreases to approximately unity for all azeotropic loadings. The number of stages does not increase linear with loading as commonly occurs for zeotropic distillation, Fair (1987). Instead F1 = 0.57 mol.min-1 appears to be an optimum loading. As azeotropic loading is increased higher solvent flowrates are required to achieve the same solvent to feed ratio, thus the packing becomes increasingly more occupied with increased xr. Bale packing has three levels of porosity. As the liquid holdup is increased a greater percentage of the liquid phase occupies the space between the small particles contained within the fibreglass quilt. This liquid cannot contribute directly to the separation, as it is not in direct contact with the vapour phase. Separation can be enhanced through either improved liquid renewal between free flowing channel liquid and the liquid within the bags or through reduced loading and lower frequency of bag occupation. As the flowrates for good renewal of the static packing are unattainable under distillation conditions, Chapter 5, then separation can only be enhanced by lowering the loading.

Chapter 6: Extractive Distillation over Bale Packing 289 Liquid phase flowrate may produce several opposing effects in an extractive distillation column over Bale packing. Meirelles and Telis (1992) demonstrated that for extractive distillation of the system ethanol/water using ethylene glycol the governing mass transfer resistance lay on the liquid film side, which is the reverse of conventional distillation. Liquid phase flowrates and mixing thus will have a very important role in column efficiency. Thus, generally for a given azeotrope loading as solvent loading or flowrate is increased efficiency improves due to lowered mass transfer resistance. However, especially for the countercurrent system of the current study at certain point the effect of entrainment of TBA downward (entrainment in reverse) prevails and recovery and efficiency drop off. Further, as previously mentioned increased loading may lead to increased bag penetration and lower effective contact area.

Chapter 6: Extractive Distillation over Bale Packing 290 1

0.9

0.8

0.7 TBA y' 0.6

0.5

0.4

0.3 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

x'TBA

equi. xr 0.46 equi. xr 0.52 equi. xr 0.62 opp. xr 0.46 opp. xr 0.52 opp. xr 0.62

-1 Figure 6.13 Analysis of runs to for Az 0.49 mol.min for experiment 1. xEG = 0.46,

2. xEG =0.52 and 3. xEG = 0.62.

Chapter 6: Extractive Distillation over Bale Packing 291 1

0.9

0.8

0.7 TBA y' 0.6

0.5

0.4

0.3 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

x'TBA

equi. xr 0.54 equi. xr 0.63 equi. xr 0.66 equi. xr 0.72 opp. xr 0.54 opp. xr 0.63 opp. xr 0.66 opp. xr 0.72

-1 Figure 6.14 Analysis of runs to for Az 0.57 mol.min for experiment 6. xEG = 0.42,

7. xEG = 0.54, 8. xEG = 0.63, 9. xEG = 0.66 and 10. xEG = 0.72.

Chapter 6: Extractive Distillation over Bale Packing 292 1

0.9

0.8

0.7 TBA y' 0.6

0.5

0.4

0.3 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

x'TBA

equi. xr 0.48 equi. xr 0.61 equi. xr 0.68 equi. xr 0.73 opp. xr 0.48 opp. xr 0.61 opp. xr 0.68 opp. xr 0.73

-1 Figure 6.15 Analysis of runs to for Az 0.71 mol.min for experiment 11. xEG = 0.37,

12. xEG = 0.48, 13. xEG = 0.61, 14. xEG = 0.68 and 15. xEG = 0.73.

Chapter 6: Extractive Distillation over Bale Packing 293 10

9 F1= 0.49 mol/min

8 F1 = 0.57 mol/min F1 = 0.71 mol/min 7

6

5 theoretical

N 4

3

2

1

0 0.3 0.4 0.5 0.6 0.7 0.8 xr

Figure 6.16 Extractive distillation Number of Theoretical stages achieved

(Ntheoretical) with variation in xr and different azeotrope feed loadings.

Chapter 6: Extractive Distillation over Bale Packing 294 6.4 Towards CED, Further Considerations Fabrication of bale or other forms of catalytic packing, operation within a laboratory type facility, as well as moderate use of chemical and power dictate column diameters in the range 40 to 50 mm and 2 to 4 m column heights. These dimensions are commonly found in published experimental studies such as that of Subawalla et al (1997) and Flato and Hoffmann (1992). The packing and column internals offer a certain range of practical vapour and liquid loading. In the case of extractive distillation a certain amount of entrainer is also required to effect azeotrope separation. The separation is also governed by mass transfer rates, which result in a particular packing efficiency. These separation requirements need to be matched with reaction requirements. In this short section investigation of the additional requirements of the reactive system of CED was initiated. The reactive system will require the additional component isobutylene, which is hydrated to TBA. It will also require a certain amount of catalyst Amberlyst 15 given intrinsic kinetics, effectiveness factor and reactor mass transfer in isobutylene in particular.

6.4-1 The Effect of Isobutylene Upon Selectivity Under atmospheric conditions the solubility of isobutylene is limited. The amount, which will dissolve, however, is likely to have an effect upon liquid phase activities. Given that isobutylene is non-polar and in particular more so than TBA, it can be predicted that this affect, although minor will be to suppress the separation using the polar solvent ethylene glycol. This follows from the fact that extractive distillation entrainers of polarities closer to TBA than that of water result in reversal of natural tendency of separation of the TBA/water system as opposed to enhancement as provided by ethylene glycol.

-05 The low solubility of isobutylene determined in Chapter 4, of the range xIB 10 to 0.01 in mixtures of water/EG/TBA over 316 to 353 K, suggested that the effect of addition of isobutylene would be well below that of experimental uncertainty and would not be determinable experimentally. Thus, its effect was estimated using the VLE techniques already detailed. The necessary binary parameters for the Wilson model could not be found and thus UNIFAC was turned to provide an initial estimate of VLE of the four

Chapter 6: Extractive Distillation over Bale Packing 295 component systems. The results were analysed in terms of a solvent and isobutylene free system or a pseudo-entrainer.

The effect of isobutylene for xr = 0.67 was demonstrated in Figure 6.17 (a). It was found that increasing isobutylene dropped the effectiveness of the entrainer. The effect was marginal up to xIB = 0.05 as may be expected. The shape of the x’-y’ curve was unaffected since a uniform entrainer composition was assumed. In the catalytic extractive column the solubility of isobutylene will vary over the length of the column, diminishing towards the top as the isobutylene of the vapour phases is diminished reducing the concentration driving force. This effect may allow for higher purities of TBA to be obtained. Selectivity was calculated in much the same manner as for Section 6.1 and reported in Figure 6.17 (b). It was found that selectivity dropped by 14 to 16 % with isobutylene at xIB = 0.1. The addition of isobutylene will require that greater amounts of ethylene glycol be used to achieve the same separation. The requirements of separation and reaction are thus opposing, as reaction rate, Chapter 4, is highly dependent upon isobutylene liquid phase concentration.

Chapter 6: Extractive Distillation over Bale Packing 296 1

0.8

0.6 TBA

y' xIB 0 0.4 xIB 0.001 xIB 0.01 0.2 xIB 0.05 xIB 0.1 0 0 0.2 0.4 0.6 0.8 1 (a) x'TBA

2 1.8 1.6 1.4 1.2 1 TBA.W S 0.8 0.6 0.4 xr 0.33 0.2 xr 0.67 0 0 0.05 0.1 0.15 (b) xIB

Figure 6.17 Effect of isobutylene upon the extractive distillation of water and TBA

with ethylene glycol (a) x’-y’ curves in TBA and (b) selectivity as a function of xIB.

Chapter 6: Extractive Distillation over Bale Packing 297 6.4-2 Matching Separation and Reaction Rates over Bale Packing This section explores the problem of matching reaction rate with separation rate. The question of interest was, can the rate of azeotrope successfully separated in this study be generated over five bales containing approximately 185 g of Amberlyst 15? Further more how many bales are required given the rates of reaction, which can be achieved, and the solubility of isobutylene, which can be obtained?

In order to begin answering these questions the bales were represented by a CSTR saturated with isobutylene. The reaction rate required to generate the experimental azeotrope flowrates was calculated at 373 K, over 185 g Amberlyst 15 for a conversion of XW = 0.66 (corresponding to the azeotrope composition). The required isobutylene based upon the kinetic model of Equation (4.) and assuming no inhibition in TBA, at various significant xr was also calculated and the results reported in Table 6.12. It was -03 -1 found that the required CIB were not unrealistic at 3 to 6 × 10 mol.L . If an effectiveness factor of 0.1 is obtained then these increase to 3 to 6 × 10-02 mol.L-1, which may still be obtainable in a TBA rich system as indicated by the isobutylene solubility study Section 4.. These relatively simple calculations indicate that given adequate supply in terms of rate and equilibrium determined maximum for isobutylene liquid phase concentration, rates of azeotrope as separated in this study should be obtainable over five bales.

Chapter 6: Extractive Distillation over Bale Packing 298

Table 6.12 Magnitude of reaction rates and isobutylene solubility required to attain the experimental flowrates of azeotrope used in this study, at 373 K, over 185 g Amberlyst 15, in an ideal CSTR.

F1 (Az) FW CSTR FW0 CSTR rCSTR mol.min-1 mol.s-1 mol.s-1 mol.g-1.s-1 0.49 2.76E-03 8.11E-03 2.89E-05 0.57 3.23E-03 9.49E-03 3.39E-05 0.71 4.02E-03 1.18E-02 4.22E-05 0.82 4.66E-03 1.37E-02 4.89E-05

xr: 0.67 0.75 0.8 CW: 1.86 1.44 1.16 CTBA: 3.62 2.79 2.26 CEG: 11.13 12.68 13.68

F1 (Az) CIB CIB CIB mol.min-1 mol.min-1 mol.min-1 mol.min-1 0.49 0.0024 0.0033 0.0044 0.57 0.0028 0.0039 0.0051 0.71 0.0034 0.0049 0.0064 0.82 0.0040 0.0057 0.0074

avg: 0.0031 0.0044 0.0058

Chapter 6: Extractive Distillation over Bale Packing 299 6.5 Concluding Remarks Both the Wilson and UNIFAC models with available parameters captured the azeotropic behaviour of TBA and water. Analysis of x’-y’ pseudo-binary curves generated using the Wilson model with the experimentally obtained parameters of Liu et al (1993) suggested that the minium required ethylene glycol fraction to break the azeotrope was xr.min = 0.063, that the minimum fraction to avoid pinching of the x’-y- curve was xrmin.non-pinch = 0.165.

Experiments conducted over five bales in a semi-batch mode demonstrated that the azeotrope between TBA and water is overcome with the use of ethylene glycol. Over the range of the study conducted at zero reflux an azeotrope loading of F1 = 0.57 mol.min-1 and a solvent to feed ratio of Fr = 1.58 mol.mol-1 or xr = 0.42 were found to be optimum.

Analysis of experimental data over five bales in terms of the Wilson model suggested that 0.6 m of Bale packing for a given azeotrope loading was effective for feed to solvent molar ratios, xr up to a certain value beyond which severe reverse entrainment -1 occurred. These values occurred at xr = 0.52 for F1 = 0.49 mol.min , xr = 0.54 for F1 =

0.57, and xr = 0.61 for F1 = 0.71. The number of theoretical stages obtained beyond these values to tend to unity, with the operating and equilibrium lines of the particular section becoming highly divergent creating a broad operating space.

Use of VLE generated data using the UNIFAC model for the four component system including TBA/water/EG/isobutylene, suggested that the addition of isobutylene lowered the selectivity of the pseudo-singular entrainer containing both ethylene glycol and isobutylene, compared to pure ethylene glycol.

In the next section additional requirements of the reactive system brought about by mass transfer effects and synergistic reaction separation effects were investigated building upon the kinetic, fluid dynamics and separative studies conducted thus far.

Chapter 6: Extractive Distillation over Bale Packing 300 6.6 Nomenclature amk UNIFAC parameter aijs relative volatility in the presence of solvent aij relative volatility

Ai Antoine Equation Coefficient ek UNIFAC parameter f fugacity -1 F1 Column azeotropic feed molar flow rate (mol.min ) -1 F2 Column solvent feed molar flow rate (mol.min ) -1 F3 Column Distillate molar flow rate (mol.min ) -1 F4 Column Bottom Stream molar flow rate (mol.min ) -1 Fi.A Molar flow of component A in stream i, (mol.min )

Fr Ratio of solvent to azeotropic feeds, (F2/F1) (Eqn (6.19))

Ji UNIFAC parameter

Li UNIFAC parameter N Number of theoretical plates P Total Pressure (mmHg) or (kPa) sat Pi Vapour Pressure of Component i (mmHg)

PTBA Purity of TBA, distillate stream q UNIFAC parameter

Qk UNIFAC parameter r UNIFAC parameter

Rk UNIFAC parameter -1 -1 rTBA reaction rate (mol.g .s )

RTBA Recovery of TBA

ReG Gas Phase Reynolds Number, ρUGdr/μ

ReL Liquid Phase Reynolds Number, ρULdr/μ sk UNIFAC parameter

Sij Selectivity of Solvent U Superficial velocity based on reactor correctional area, (m/s) -1 Vi Molar volume (mol.kg ) xr molar fraction based on solvent to azeotropic feed ratio (Eqn (6.20)) xi liquid phase molar fraction of component i

Chapter 6: Extractive Distillation over Bale Packing 301 x’i solvent free liquid phase mole fraction yi vapour phase molar fraction of component i y’i solvent free gas phase mole fraction

Greek:

βik UNIFAC parameter γ activity coefficient μ viscosity (N.s.m-2)

θk UNIFAC parameter ρ density (kg.m-3)

τmk UNIFAC parameter Φ ratio of fugacities

Λij Wilson parameter

Acronyms: CED Catalytic Extractive Distillation CFBR Countercurrent Fixed Bed Reactor EG Ethylene Glycol HETP Height Equivalent to Theoretical Plate IB Isobutylene RTD Residence Time Distribution TBA tert-Butyl Alcohol

Chapter 6: Extractive Distillation over Bale Packing 302 6.7 Literature Cited Atkins, G. T. and C. M. Boyer “Application of McCabe-Thiele method to extractive distillation calculations.” Chem. Eng. and Progress 45(9): 553-562. (1949). Fair, J. R. (1987). Distillation. Handbook of Separation Process Technology. R. W. Rousseau. Brisbane, John Wiley & Sons: 229-239. Flato, J. and U. Hoffmann “Development and Start-up of a Fixed Bed Reaction Column for Manufacturing Antiknock Enhancer MTBE.” Chem. Eng. Tech. 15: 193-201. (1992a). Gmehling, J. and U. Onken "Vapour Liquid Equilibrium Data Collection, Aqueous Organic Systems". Frankfurt, DECHEMA. (1977). Herington, J. J. Inst. Petr. 37: 457. (1951). Kumar, S., J. D. Wright and P. A. Taylor “Modelling and Dynamics of an Extractive Distillation Column.” Can. J. Chem. Eng. 62: 780-789. (1984). Laroche, L., N. Bekiaris, W. Anderson and M. Morari “The Curious Behaviour of Homogeneous Azeotropic Distillation-Implications for Entrainer Selection.” AIChE Journal 38(9): 1309-1328. (1992b). Liu, F., C. Zhang, F. Huang and C. Zhang “Studies on separation of alcohols and water by extractive distillation.” Fuel Science and Technology International 11(11): 1537-1550. (1993). Lynn, S. and D. N. Hanson “Multieffect extractive distillation for separating aqueous azeotropes.” Ind. Eng. Chem. Process Des. Dev. 25: 936-941. (1986). Meirelles, A., S. Weiss and H. Herfurth “Ethanol Dehydration by Extractive Distillation.” J. Chem. Tech. & Biotechnology 53: 181-188. (1992c). Momoh, S. O. “Assessing the accuracy of selectivity as a basis for solvent screening in extractive distillation processes.” Separation Science and Technology 26(5): 729-742. (1991). Oyer, R. V. and J. M. Prausnitz Ind. Eng. Chem. 57(5): 18-26. (1965). Pretel, E. J., P. A. Lopez, S. B. Bottini and E. A. Brignole “Computer-Aided Molecular Design of Solvents for Separation Processes.” AIChE Journal 40(8): 1349-1360. (1994). Resa, J. M., M. A. Betolaza, C. Gonzalez and A. Ruiz “Isobaric vapor-liquid equilibria of acetone-propyl ether and isopropyl ether systems. Corroboration of no Reverse volatility.” Fluid Phase Equilibria 110: 205-217. (1995).

Chapter 6: Extractive Distillation over Bale Packing 303 Resa, J. M. and G. A. Ruiz “Experiments of extractive distillation at laboratory scale for the rupture of the azeotropic mixture acetone and isopropyl ether.” Separation and Purification Technology 18: 103-110. (2000). Ruiz, C., J. Coca, A. Vega and F. V. Diez “Extractive distillation of hydrocarbons with dimethylformamide: Experimental and simulation data.” Ind. Eng. Chem. Res. 36: 4934-4939. (1997a). Sinnott, R. K. "Coulson and Richardson's Chemical Engineering". Sydney, Butterworth and Heinemann. (1997b). Smith, J. M., H. C. V. Ness and M. M. Abbott "Introduction to Chemical Engineering Thermodynamics", McGraw-Hill International Editions. (1996). Stichlmair, J. G. and J. R. Herguijuela “Separation Rgions and Processes of Zeotropic and Azeotropic Ternary Distillation.” AIChE Journal 38(10): 1523-1535. (1992d). Subawalla, H., J. C. Gonzalez, A. F. Seibert and F. J. R. “Capacity and Efficiency of Reactive Distillation Bale Packing: Modeling and Experimental Validation.” Ind. Eng. Chem. Res. 36: 3821-3832. (1997c). Wilson, G. M. “Vapour-Liquid Equilibrium XI A new expression for the excess free energy of mixing.” J. Am. Chem. Soc. 86: 127-130. (1964). Yang, B. and H. Wang “Vapour Liquid Equilibrium for Mixtures of water, alcohols and ethers.” J. Chem. Eng. Data 47: 1324-1329. (2002).

Chapter 6: Extractive Distillation over Bale Packing 304 7. Synthesis of TBA under CFBR and CED Conditions in the Presence of Ethylene Glycol

This Chapter deals with the synthesis of TBA by isobutylene hydration over Amberlyst 15 in the presence of the solvent ethylene glycol over bale catalytic packing within two pilot scale three-phase reactors. The kinetics of this reaction were investigated in Chapter 4. The Amberlyst 15 in this study was contained within the fibreglass bags of bale packing. The bales allowed for countercurrent gas and liquid operations over a wide range of flowrates, without flooding or excessive pressure drop. The fluid dynamics and separation properties of bale packing were experimentally determined in previous Chapters 5 and 6. The countercurrent arrangement being made use of the synthesis of TBA was carried out in a Countercurrent Fixed Bed Reactor (CFBR) and in a Catalytic Extractive Distillation (CED) Column.

Under the CFBR arrangement water and ethylene glycol were fed as a liquid at the top of a bed of catalytic bale packing containing Amberlyst 15. The gas phase was fed countercurrently at the base of the bed and was composed of isobutylene and nitrogen. The only separative action of this arrangement was the retention of liquid reagents in the liquid phase and gaseous reagents in the gas phase. The product TBA had little to no means of separation and exited the reactor in the same stream as any undesirable condensed product such as MET. Using this arrangement, several aspects of the proposed process of CED using ethylene glycol could be tested. These aspects included the mass transfer of isobutylene under simplified conditions, improvement in isobutylene solubility and the effect of mass transfer effects upon the selectivity ratio. Determination of the governing mass transport over bale packing was the aim of the CFBR study.

Under CED conditions ethylene glycol was fed at the top of the column and the reagents were both fed in gaseous form at the base. Steam and isobutylene were absorbed by the ethylene glycol, which functioned as a medium for liquid phase reaction. The TBA

Chapter 7: Synthesis of TBA... 305 formed was distilled off with excess steam and formed the top distillate. The promotional effect of the entrainer was investigated. The aim of this study was to identify the path of TBA and MET through the process and ascertain their extent and rate of formation over bale packing.

This chapter first describes relevant mass transfer correlations as earlier introduced in the literature review of Chapter 2. The CFBR and CED studies are then discussed in two separate sections.

7.1 Three Phase Bale Reactor Mass Transfer In order to interpret any catalytic study on a pilot scale a good understanding of the underlying physical properties, kinetic, fluid dynamics and mass transfer are required. Thus far mass transfer has only been dealt with in terms of overall observable effects present in the extractive distillation separation study. However specific mass transfer steps pertaining to a relevant model are required, especially for the transfer of the limiting reagent isobutylene.

A relevant form of analysis for differential reactors as a simple model is that of the mass transfer resistance model or formulation of an overall rate equation. For three phase reactors such as slurry reactors, trickle bed reactors and CFBRs the reactants in the gas phase participate in the following steps:

1. Diffusion from the bulk gas phase to the gas-liquid interface 2. Diffusion in the liquid phase from the interface to the bulk liquid 3. Diffusion from the bulk liquid to the external surface of the catalyst 4. Internal diffusion within the porous catalyst 5. Reaction within the porous catalyst

While the external mass transfer steps are treated separately to each other, internal mass transfer and reaction are lumped and expressed in terms of effectiveness η. The standard expression, obtained through the addition of the steps for a trickle bed reactor and pseudo first order reaction kinetics, takes the form, Fogler:

Chapter 7: Synthesis of TBA... 306 CAi −=rA (7.01) ⎛⎞()11−−ερ() ερ 11 ⎜⎟cc+++ ⎜⎟′ ⎝⎠HkGi a k Li a k s a p η k

where (1-ε)ρc is the loading of catalyst. In a trickle bed reactor ε is simply the porosity between the packed particles. For a CFBR, this term should also include the packing voidage and fraction of the reactor occupied by inert packing. The path of mass transport will be affected by this inert structure, however, through correlation of experimental data the overall influence may be characterised. Additionally, the interfacial and effective areas also require experimentally derived correlations.

Zheng and Xu (1992) have previously characterised the mass transfer over Bale packing for the three external steps mentioned. Their model was specifically developed for bale type catalytic packing based upon data from a 45.8 mm glass column. They employed bales of height h = 200 mm filled with Amberlyst 15 of 300 to 1200 μm, of at 1545.6 m2.m-3, with a bed height of 0.6 m. They worked with an aqueous system using temperature-controlled conditions. Their experimental conditions were in all respects consistent with those of this study. Employing the film theory for bulk gas to liquid transfer steps, they obtained: 1. Gas film mass transfer coefficient:

⎛⎞kadGp −3 0.92 0.24 0.5 ShGaGaLG==×⎜⎟ RT1.072 10 Re,, Re Sc (7.02) ⎝⎠aDtG

2. Liquid film mass transfer coefficient:

⎛⎞kad Sh==Lp 0.149Re0.30 Sc 0.5 L ⎜⎟ aL, L (7.03) ⎝⎠aDtL

3. Liquid solid mass transfer coefficient:

Chapter 7: Synthesis of TBA... 307 ⎛⎞2/3 kas ⎛⎞μL −−0.27 0.28 J Da==⎜⎟⎜⎟0.586Re,,G ReaL ⎜⎟aUρ D (7.04) ⎝⎠tL⎝⎠ LL where:

ae 0.24 = 0.053ReaL, (7.05) at and

4εbale d p = (7.06) at

For Equations (7.02) to (7.05) Reynolds number was defined as:

4Uiiρ Reai, = (7.07) atiμ

Chapter 7: Synthesis of TBA... 308 7.2 Synthesis of TBA over Bale Packing Operated under CFBR Conditions Two types of CFBR runs were performed, single pass and recirculation runs. The objective of the single pass runs was to gauge the extent of conversion and obtain readily characterised conditions of minimal product formation. Using the mass transfer correlations of Zheng and Xu (1992), the resistances to reaction were estimated and the reaction runs analysed. The objective of the recirculation study was to allow moderate product accumulation, such that the relative rate of TBA and MET formation could be obtained. Detailed in this section are the experimental procedure, experimental design, calculation procedure, results and analysis of the CFRB runs.

7.2-1 CFBR Experimental Described in the following section is the experimental CFBR procedure and its development, the variables chosen for investigation and the range selected. The variables considered to be of interest were gas and liquid flow rate, ethylene glycol concentration and gas feed isobutylene molar fraction.

7.2-1-1 Analytical Procedure The analysis of TBA in water/EG and of water in TBA has been previously described in Sections 4.2-3 and 6.2-2 respectively. In this section the analytical procedure used for determination of the gas phase concentration of isobutylene in nitrogen is summarised. A Shimadzu GC-8A TCD was used equipped with an Alltech packed column, BMEA o 20 wt % on C/Sorb, ss 1/8”, 24’, and with Tmax 150 C. The GC and Integrator settings are reported in Table 7.01. An injection loop volume of 1 mL connected to a six way Valco valve was used downstream from a flow regulated process stream split system allowing a sample flow of 10 mL.min-1. The following correlation for isobutylene concentration was obtained:

−08 −03 CAIB =×1.6645 10IB +× 1.0231 10

Chapter 7: Synthesis of TBA... 309 Table 7.01 GC and Integrator Settings for gas phase analysis of isobutylene GC-8A TCD Gas Analysis Variable Value Variable Value Carrier 1, He 240 kPa Attenuation 4 Carrier 2, He 80 kPa Stop time 6 Inj/Det Temp. 90 oC Slope 300 Col. Temp. 80 oC Width 5 Current 70 mA Minimum Area 10 Polarity Plus Attenuation 1

7.2-1-2 CFBR Procedure Using the rig of Figure 7.01, the hydration of isobutylene was performed in the CFBR for both runs containing the proposed solvent ethylene glycol and ones with water alone. Single pass and recirculation type runs were performed. TBA and MET were analysed in the effluent streams and within the recirculation tank, by GC-FID as detailed in Section 4.2-3. The unreacted isobutylene was also determined in the exit stream, Section 7.2-1-1. No products (TBA or MET) were detected in the gas outlet stream suggesting that the condenser used was effective at retaining condensable components within the system.

All experiments were performed under isothermal conditions at 348 K. The feed was preheated before entering the reactor by passing it through a 16 m loop immersed in a temperature controlled bath set at 358K. The temperature of the preheater bath was marginally adjusted depending on flowrate and ambient conditions. The jacket of the column was heated with a silicon oil recirculation bath set at 368K. The system was insulated with three layers of specialist fibreglass insulation tape as used on the steam supply lines. The average reactor temperature based upon three measurement points along the bed was 349 ± 0.25 K.

The reactive section was packed with five bales containing the catalyst Amberlyst 15. Flowrates were monitored and determined through the use of rotameters calibrated at the temperature of operation. The liquid level within the liquid collector was adjusted manually through the use of a manual valve on the liquid outlet line.

Chapter 7: Synthesis of TBA... 310

P [03] T [01] P T T [02]

[04]

T T1 [06]

[05] T2

[07] T3

T

T4 [09] [10]

[08]

[11]

[12]

P P

[15] [13]

[14]

Figure 7.01 CFBR rig for isobutylene hydration over Bale Packing containing Amberlyst 15

Key to Figure 7.01 CFBR rig for isobutylene hydration over Bale Packing containing Amberlyst 15 [01] Temperature Indicator and Thermocouple Selector Switch, Shinko Controller, RS Selector.

Chapter 7: Synthesis of TBA... 311 [02] Condenser, Single Coil Unjacketed, QuickFit 19/26, Crown Scientific. [03] Gas Sampling Assembly, 6 Way Valco Valve, Whitey Needle Valves, Soap bubble meter 100 mL. [04] GC-8A TCD, Shimadzu. [05] Liquid Feed Preheater: Silicon oil bath with Braun Bath Heater, SS bath, Heat Exchanger Coil, ¼” SS 40 Rectangular Loops at 40 cm/loop. [06] Reactor Head: Liquid distributor, Spray arrester (5 cm of 8mm Raschig Rings), Gas exhaust line. [07] Reactor Column: Jacketed Column, 45 mm id, 60 mm od. [08] Stirrer: Drive and Controller, Hargil Dynamics. [09] Liquid Feed Rotameter: Planton, Duff and Machintosh, Model 48, universal, with scale marked for 2 to 25 L/min Air. [10] Jacket Heating Oil Bath: Bath Heater and Recirculator, Haake DC 5; Insulated SS tank. [11] Liquid Feed Pump: Masterflex Peristaltic Pump, Variable Speed Variable Speed Modular Drive, Model 7016 pump head, L/S 16 Viton Tubing, Cole Parmer. [12] Reactor Base: Liquid Receiver and level sight, gas and liquid connections, built inhouse of copper 50 mm tube. [13] Liquid Feed Tank: SS 15 L. [14] Liquid Waste Tank: 20 L polypropylene waste container. [15] Gas Supply: Built inhouse in a stable/sturdy aluminium frame. Includes: 3 × 10” Rotameters: Fisher Porter Meter Tube No FP ¼-20-G-5/84 Meter Tube No FP 1/8-12-G-5/84 Meter Tube No FP 1/8-08-G-5/84 1 × 6” Rotameter: Fisher Porter Meter Tube No FP 1/8-08-P-3/37 2 × Gas Regulator: Fairchild Regulator, Kendall, Model 10, Range 5 to 400 PSI 2 × Pressure Gauges: Dobbie Instrument Australia 0 to 250 kPa 2 × Needle Valves, Whitey SS ¼”

Chapter 7: Synthesis of TBA... 312

Stepwise CFBR Recirculation Experimental Procedure The following are step by step instructions for the CFBR recirculation experiments. The instructions for single pass experiments were similar, instead of recirculation the outlet of the reactor was drained to waste.

1. Prepare GC-8A FID, (need at least 2 hr for this step). 2. Prepare GC-8A TCD for gas analysis. 3. Prepare Tank Charge. 4. Turn on condenser water, 2 to 3 L/min. 5. Turn on Preheater bath [05] and Jacket bath [10]. 6. Turn on pump adjust flow rate for Reactor warm up. 7. Ensure the reactor return line is clear, set three way Liquid Feed bypass valve to column and bottom bypass valve to Feed tank [13]. 8. Allow time for temperature stabilisation; adjust tank temperatures as necessary to achieve desired experimental temperature. 9. Prepare gas feed [15], turn three way gas valves to allow sampling of feed and bypass of reactor. Use the smaller bubble meter [03] and adjust the sample flowrate to 10 mL /10 s, 60 mL/min. 10. Allow gas sample flow to stabilise and sample by turning Valco Valve to sample [03]. Sample feed at least twice, check for reproducbility, areas within 10%. 11. Turn gas three way valves to reactor and start timer. 12. Take Feed tank 3 mL samples every 20 min, filter with 0.22 μm filter, and inject on GC-8A FID. 13. Adjust gas sample flow, sample gas every 20 min, starting at 30 min. 14. Monitor all flow meters during the experiment, cycle between them, record the temperature every 30 min. 15. Shut down: turn off gas, turn off liquid feed, empty feed tank, wash reactor with a 10 % solution of ethylene glycol in water or distilled water alone. Turn off heating baths.

Chapter 7: Synthesis of TBA... 313 7.2-1-3 Experimental Design The design of the single pass study was relatively simple. The feed to the CFBR, hence

ReL, would be increased until the limit of detection of both possible products TBA and

MET was reached. The gas flowrate would be varied to ascertain the influence of ReG. The temperature was maintained at 348 K, well below the boiling point of TBA, yet sufficiently high for good reaction rates to be obtained. The aim of these runs was to evaluate the extent of conversion of single pass reaction over five bales. To analyse the mass transfer of isobutylene to the reaction sites in terms of available mass transfer correlations.

The recirculation study was conducted in order to gain a better understanding of the selectivity towards TBA for hydration in the presence of ethylene glycol. Recirculation would allow for sufficient MET to build up allowing for rate determination. The variables chosen for the study were:

A. Isobutylene Gas Phase Molar Fraction. The gas phase molar fraction of isobutylene will drop across an integral large scale CFBR, thus, for the differential reactor of these experiments, the feed isobutylene concentration was varied to determine the effect of decreased concentration driving force upon reaction rate and

selectivity. The range of yIB chosen was 0.3 to 0.6 diluted in nitrogen. This allowed for determinable concentrations and covered the range of isobutylene as may be obtained by catalytic cracking, Jayadeokar and Sharma (1993). -1 B. Liquid Flowrate, QL (L.min ). The flowrate of liquid directly influences a degree of backmixing, isobutylene mass transfer and liquid retention. The importance of this variable for CFBR reaction was demonstrated in the single pass runs. The range -1 was chosen between 0.1 and 0.25 L.min giving an ReL range of 20 to 80. C. Ethylene Glycol Liquid Molar Fraction. The range chosen was 0.3 to 0.6. The range is of interest for separation as demonstrated in Chapter 6. This range also coincided with a range of selectivity factor crossover as determined within the

kinetic study of Chapter 4, Equation (4.64 & 4.65). At xEG of 0.4 the selectivity factor dropped below 1. Variables, which were fixed for the purpose of the study, were catalyst loading, temperature, and gas flowrate. Gas phase mixing was found to be more sensitive to ReL

Chapter 7: Synthesis of TBA... 314 than ReG thus gas flowrate was kept constant. The variables of temperature and catalyst loading should be investigated once the CED design for TBA synthesis is better established.

In order to probe the effect of isobutylene and ethylene glycol molar fractions and liquid flowrate, a full factorial centre point experimental design was employed. Experimental Design and the centre point design are well detailed in Montgomery (1997). The centre point design is a full factorial design which, incorporates an additional, nc, centre points. The set point of these are the averages of the high and low levels (-1,1) of the factors and are denoted by "0”. The replication of the centre point acts as an estimate of the error of the measures and allows for an estimate of the significance of curvature across the range of the levels, SSpq (pq: pure quadratic). This is useful when non-linearity is a possibility. Experimental design was significantly utilised throughout this chapter, for the following reasons:

• The current pilot scale experiments were long and intensive with considerable analysis time, especially for runs employing ethylene glycol. Further, they consumed considerable chemicals and resources. The scale was necessary for packing fabrication, relevance of the study, as well as comparison and use of established correlations. Experimental design provided an opportunity for valuable information to be collected economically, while defining the region progressively. • There was a high chance of interaction of the variables, which, using a matrix type design, as used for correlation development in previous chapters, would require a large volume of experiments to be determined. • The possibility of little observable reaction rate suggested that segmented bounds be evaluated first and dependent on features of particular interest filled in at a later stage. • In both the case of the CFBR and CED study, the experiments were an indication of the final process or design direction path, and not a complete representation of an existing process, which may require characterisation for the purpose of commercial optimisation. The factors were varied according to Table 7.02. Each of the three factors was varied 3 over two levels, giving a full factorial, experimental design of nf = 2 = 8. Three centre

Chapter 7: Synthesis of TBA... 315 point experiments, ncp = 3, were conducted in order to determine the curvature of the measures and estimate the error of experimentation.

Table 7.02 Variable Ranges for the CFBR Recirculation Study Variable Variable Low Level Centre Point High Level A yIB 0.3 0.45 0.6 B QL 0.101 0.162 0.223 C xEG 0.3 0.45 0.6

The measures of interest were the reaction rates of TBA and MET formation, the overall effectiveness of these reaction rates when compared to the kinetic study and their ratio the selectivity factor. 7.2-1-4 Calculation Procedure For single pass experiments under isothermal conditions a balance in terms of the product, A, over the reactor and total condenser gave:

dVC A =+−QC mrQ C (7.08) dt Lin,, Ain A Lout,, Aout

Given steady state conditions with no accumulation and Ci0 = 0 this reduced to:

−−11 QCL,, out A out rmolgsA ().. = (7.09) m and was used to calculate product reaction rates. Several samples, taken at regular intervals, allowed for steady state to be validated and the steady state average was then used in calculation.

For recirculation experiments the bale packed CFBR was assumed to be a differential reactor. This assumption is valid given the low conversion in isobutylene and water observed for single pass experiments. The standard differential reactor procedure was employed as commonly applied to conventional trickle bed reactor analysis, Herskowitz (1979) and Morita and Smith (1978). The reaction rate was followed through the change of product concentration within the recirculator. The batch system was assumed to be

Chapter 7: Synthesis of TBA... 316 closed, thus QL,in and QL,out = 0 and 100 % efficiency for the condenser was assumed. A balance of the form of Equation (7.08), conducted over the whole system reduced to:

−−11 Vtot ⎛⎞dCA rmolgsA ().. = ⎜⎟ (7.10) mdt⎝⎠

where Vtot was the volume of the liquid within the system. The bed was completely drained after washing with distilled water and process lines drained with Vtot being the volume of the initial charge to the recirculation tank [13]. This volume of 3 to 4 L was much greater than that of the reactor liquid holdup HL0 ≈ 0.15 L. The concentration was followed with time and the slope of the resulting plot employed to complete the calculation of rA. For the experimental design of recirculation study the measures of interest were: rTBA, rMET, overall effectiveness factor η0 and selectivity factor ξ. The overall effectiveness was defined as:

rA,observed η0 = (7.11) rA,intrinsic

where rA,intrinsic had been determined previously in Chapter 4 in a rotating basket reactor. The selectivity factor was defined as the ratio:

r ξ = TBA, observed (7.12) rMET, observed

The standard ANOVA procedure outlined by Montgomery (1997) was followed. The Contrast for each variable and interaction was calculated according to:

Contrastijj= ∑ν y (7.13) where ν are either low or high levels of the design table –1 or +1. The Effect is defined as: Contrast Effect = (7.14) n2k −1 The sum of squares of the variables are given by:

Chapter 7: Synthesis of TBA... 317

()Contrast 2 SS = (7.15) n2k and the mean square by:

SSi MSi = (7.16) dofi

For a centre point design the sum of square of the error SSe and of the curvature SSpq are given by:

n cp 2 SSecpicp=−∑() y, y (7.17) i=1

2 nnfcpf( y− y cp) SS pq = (7.18) nnfcp−

The mean square of the error was given by:

SSe MSe = (7.19) ncp −1

The F statistic was calculated using:

MSi Fi = (7.20) MSe

The F probability distribution, Pi, between a factor and the error was calculated using the function, FDIST, within Microsoft Excel, with syntax of:

Pii= FDIST( F,, dofi dofe) (7.21)

Chapter 7: Synthesis of TBA... 318

7.2-2 Results of Single Pass CFBR Runs The single pass CFBR runs were conducted at steady state, which was determined through sampling every 20 minutes with subsequent analysis. The raw TBA concentrations were reported in Figures 7.02 (a) for water runs with ReG 22 and (b) ReG 11. Figure 7.03 depicted the raw results of runs conducted with 30 mol % ethylene glycol. The ethylene glycol containing runs tended to give a higher degree of scatter due to the higher error of the analysis in the presence of solvent. MET formation was not observed during the 30 mol % runs. It is likely that the amount of MET formed was below the limit of detection of MET of 0.002 mol.L-1. The desire to evaluate the selectivity factor, under mass transfer limited conditions, while starting with CA0 = 0 for each product, promoted the use of recirculation for further experimentation.

A summary of the calculated results was presented in Table 7.03. The concentration of the product for runs of higher ReL began to approach the limit of detection of TBA.

Table 7.03 Results of the Single Pass Water and Ethylene Glycol Containing Runs

Run yIB QG xW QL CTBA avg. rTBA (L.s-1) (L.s-1) (mol.L-1) mol.g-1.s-1 W1 0.6 1.667E-02 1 5.530E-03 0.0009 2.634E-08 W2 0.6 1.667E-02 1 4.330E-03 0.0013 2.985E-08 W3 0.6 1.667E-02 1 3.129E-03 0.0012 1.989E-08 W4 0.6 1.667E-02 1 1.929E-03 0.0023 2.351E-08 W5 0.6 1.667E-02 1 7.281E-04 0.0051 1.991E-08 W6 0.6 8.333E-03 1 5.530E-03 0.0010 2.862E-08 W7 0.6 8.333E-03 1 4.330E-03 0.0011 2.522E-08 W8 0.6 8.333E-03 1 3.129E-03 0.0014 2.389E-08 EG1 0.6 1.667E-02 0.7 3.036E-03 0.0029 4.725E-08 EG2 0.6 1.667E-02 0.7 2.357E-03 0.0043 5.492E-08 EG3 0.6 1.667E-02 0.7 1.677E-03 0.0056 5.109E-08 EG4 0.6 1.667E-02 0.7 9.980E-04 0.0050 2.684E-08 EG5 0.6 1.667E-02 0.7 3.185E-04 0.0092 1.591E-08

Chapter 7: Synthesis of TBA... 319 0.006 ReL 399 0.005 ReL 312 ReL 226

) 0.004 -1 ReL 139 ReL 53 0.003 (mol.L TBA

C 0.002

0.001

0 02468(a) Sample (1/20 min-1)

0.002 ReL 399 0.0018 ReL 312 0.0016 ReL 226

) 0.0014 -1 0.0012 0.001 (mol.L 0.0008 TBA

C 0.0006 0.0004 0.0002 0 02468 (b) Sample (1/20 min-1)

Figure 7.02 CFBR liquid outlet concentration of product TBA for Single Pass

Water Runs (a) ReG = 22 and (b) ReG = 11.

Chapter 7: Synthesis of TBA... 320 0.014 ReL 65

0.012 ReL 50 ReL 36 0.01 ReL 21 )

-1 ReL 7 0.008

(mol.L 0.006 TBA C 0.004

0.002

0 012345 Sample (1/20 min-1)

Figure 7.03 CFBR liquid outlet concentration of product TBA for Single Pass runs using 30 mol % ethylene glycol.

Chapter 7: Synthesis of TBA... 321 7.2-3 Discussion of Single Pass Study As shown in Figure 7.04 (a) the reaction rate of TBA formation of the Single Pass study was directly influenced by ReL, while ReG over the range tested had little measurable influence. The strong influence of ReL suggested external mass transfer limitation.

3.5E-08

3.0E-08 )

-1 2.5E-08 .s

-1 2.0E-08 1.5E-08 (mol.g 1.0E-08

TBA ReG 22 r 5.0E-09 ReG 11 0.0E+00 0 100 200 300 400 500 (a) ReL

8.E-08 7.E-08 )

-1 6.E-08 .s

-1 5.E-08 4.E-08 3.E-08 (mol.g 2.E-08 TBA r 1.E-08 0.E+00 020406080 (b) ReL

Figure 7.04 Observed TBA Reaction for (a) solvent absent water runs and (b) solvent present, ethylene glycol runs.

Chapter 7: Synthesis of TBA... 322 The observed reaction rate was higher in the presence of ethylene glycol, Figure 7.04 (b). The higher reaction rates could be attributed to improved isobutylene solubility.

The reaction rate in the presence of ethylene glycol also increased with ReL indicating that the influence of mass transfer was also limiting as in the case of the pure water runs. While the flowrates used were similar to that of the pure water single runs the high viscosity of ethylene glycol decreases the achievable ReL.

The mass transfer coefficients for the single pass runs were estimated from the correlation of Zheng and Xu (1992) (cf. Eqn 7.02 – 7.04). The required physical properties and relations were given in Table 7.04.

Table 7.04 Physical Properties and Relations used to calculate mass transfer coefficients. Property Component Value T (K) 348 density (kg.m-3) water 974.19 EG 1079.46 N2 (g) 0.96 viscosity water 3.821E-04 EG 3.696E-03 N2 (g) 2.000E-05 Diffusivity IB (m2.s-1) water 2.940E-09 EG 5.547E-10 N2 (g) 1.340E-05 Column and Packing Column Area (m2) 0.00 d (m) 0.05 2 -3 at (m .m ) 1546.00 porosity 0.71

dp (m) 1.837E-03 total mcat (g) 185

These mass transfer coefficients were employed in the overall rate model given by Equation (7.01). The assumption was made that the extent of TBA formation was small and that the rate of transport of the product away from the reactive sites would not be limiting. Further, that the observed reaction rate was limited predominantly by isobutylene transport. The resistances obtained were reported in Table 7.05. These resistances were used to estimate the observed reaction rate. The intrinsic, observed and predicted rate of TBA formation were also reported. It can be seen that the observed rates were an order of magnitude slower than the intrinsic rate. The model reaction rates

Chapter 7: Synthesis of TBA... 323 reflect this drop in rate and also follow the same trend in ReL, which obviously affects mass transfer. A parity plot of the model overall rate to that observed is presented in

Figure 7.05. The level of agreement was moderate and best at low ReL. The discrepancy may be due to experimental error of the CFBR runs. The influence of surface tension, especially upon effective area, has not been introduced into the correlation of Zheng and Xu. The agreement for water runs was marginally better than that with the ethylene glycol runs most likely due to the fact that the mass transfer correlation were developed within an aqueous system. Ethylene Glycol, compared to water, has a much higher viscosity and different wetting properties such as surface tension and contact angel. As the mass transfer over bale packing takes place largely under conditions of film flow along bag and mesh walls, the effective area for mass transfer would be highly affected by these properties. Correlations may need to be developed which are solvent specific. Further, the correlations of Zheng and Xu were most probably developed assuming plug flow conditions and may need to be corrected for dispersion. Currently, the theory of such correction is being developed and has been applied to distillation height of transfer units estimates, Marcias-Salinas and Fair (2002b).

The predicted resistances were plotted against ReL in Figure 7.06 (a) and (b). The resistances observed indicate that for pure water runs the dominant resistance is that of R2 of interface to bulk liquid. In the presence of ethylene glycol, the dominant resistance is that of R3 bulk liquid to solid at low ReL and R2 interface to bulk liquid at higher ReL. For the same ReL the external mass transfer resistance (R1+R2+R3) in the presence of ethylene glycol is lower than that of pure water.

The two significant external resistances, R2 and R3, were plotted against reactor Peclet number, Per,L,B, Figure 7.07. Reactor Peclet number was predicted by the correlation of Equation (5.63). Whilst an unorthodox plot, it serves to demonstrate backmixing effects.

The exponential line of fit was extrapolated back to Per,L,B ≈ 6 which was the optimum backmixing achieved at ReL 623. For pure water runs within the pre-loading regime it seemed unlikely that external mass transfer resistances could be eliminated with increased ReL and decreased Per,L,B. However, it may be possible, through the use of ethylene glycol, to reduce or eliminate the external mass transfer limitation. Higher degrees of backmixing, as indicated by lower Per,L,B resulted from better convective

Chapter 7: Synthesis of TBA... 324 penetration of the catalyst containing bags. However, while ethylene glycol lowers external resistances, that of the internal mass transfer and catalytic reaction are increased. Zhang, Adesina and Wainwright (2003) observed that for solvent free isobutylene hydration in a basket type, CFBR operated within the flooding regime, internal mass transfer dominated the observed reaction rate. This suggests that for a given ReG, as ReL is increased and loading and flooding points are attained, the efficiency of external mass transfer in the absence of a solvent may improve to the point that internal mass transfer becomes the dominant resistance.

Chapter 7: Synthesis of TBA... 325

Table 7.05 Mass Transfer Coefficients (kia) and Mass Transfer Step Resistances (Ri) Calculated

Run kGa kLa ksa R1 R2 R3 R4 rTBA int. rTBA obs. rTBA model s-1 s-1 g.s.L-1 g.s.L-1 g.s.L-1 g.s.L-1 mol.g-1.s-1 mol.g-1.s-1 mol.g-1.s-1 W1 1.41E-05 1.09E-02 4.67E-02 344 21350 4981 6583 4.66E-07 3.80E-08 2.63E-08 W2 1.33E-05 1.01E-02 3.92E-02 364 22976 5940 6583 4.66E-07 3.53E-08 2.99E-08 W3 1.23E-05 9.19E-03 3.10E-02 394 25327 7505 6583 4.66E-07 3.18E-08 1.99E-08 W4 1.09E-05 7.94E-03 2.19E-02 442 29285 10634 6583 4.66E-07 2.69E-08 2.35E-08 W5 8.66E-06 5.93E-03 1.08E-02 559 39224 21442 6583 4.66E-07 1.87E-08 1.99E-08 W6 7.45E-06 1.09E-02 5.63E-02 650 21350 4131 6583 4.66E-07 3.87E-08 2.86E-08 W7 7.02E-06 1.01E-02 4.72E-02 689 22976 4926 6583 4.66E-07 3.60E-08 2.52E-08 W8 6.50E-06 9.19E-03 3.74E-02 745 25327 6224 6583 4.66E-07 3.25E-08 2.39E-08 EG1 9.11E-06 9.41E-03 1.43E-02 882 24720 16261 17853 2.85E-07 3.51E-08 4.72E-08 EG2 8.57E-06 8.72E-03 1.19E-02 937 26672 19514 17853 2.85E-07 3.23E-08 5.49E-08 EG3 7.90E-06 7.88E-03 9.33E-03 1016 29537 24928 17853 2.85E-07 2.86E-08 5.11E-08 EG4 6.98E-06 6.74E-03 6.42E-03 1151 34516 36229 17853 2.85E-07 2.34E-08 2.68E-08 EG5 5.30E-06 4.78E-03 2.82E-03 1514 48619 82445 17853 2.85E-07 1.39E-08 1.59E-08

Chapter 7: Synthesis of TBA... 326

6.E-08 Water

) 5.E-08 30 mol% EG -1 .s -1 4.E-08

3.E-08

2.E-08 model (mol.g TBA r 1.E-08

0.E+00 0 1E-08 2E-08 3E-08 4E-08 5E-08 6E-08

-1 -1 rTBA,observed (mol.g .s )

Figure 7.05 Overall reaction rate observed and predicted.

Chapter 7: Synthesis of TBA... 327

45000 R1: G to i 40000 R2: i to L 35000 R3: L to s 30000 R4: cat ) -1 25000 20000 R (g.s.L 15000 10000 5000 0 0 100 200 300 400 500 (a) ReL

90000 R1: G to i 80000 R2: i to L 70000 R3: L to s 60000 R4: cat ) -1 50000

40000 R (g.s.L 30000

20000

10000

0 0 20406080 (b) ReL

Figure 7.06 Mass Transfer Resistances to Isobutylene hydration over Amberlyst 15

as a function of ReL for (a) water and (b) ethylene glycol containing systems

Chapter 7: Synthesis of TBA... 328 45000 R2: i to L 40000 R3: L to s 35000 30000 ) -1 25000 20000 R (g.s.L 15000 10000 5000 0 0 5 10 15 20 25 (a) Per,L,B

90000 80000 R2: i to L 70000 R3: L to s 60000 ) -1 50000 40000 R (g.s.L 30000 20000 10000 0 0 5 10 15 20 25 (b) Per,L,B

Figure 7.07 Mass Transfer Resistances to Isobutylene hydration over Amberlyst 15

as a function of Per,L,B for (a) water and (b) ethylene glycol containing systems

Chapter 7: Synthesis of TBA... 329 7.2-4 Results of CFBR Recirculation Runs The set points of the factors and the results were presented in Table 7.06. TBA and MET were produced in measurable quantities. Whilst the concentration of isobutylene within the gas phase was measured, these results are not reported as there was little observable change in concentration of the gas inlet to the gas outlet of the reactor. Under the conditions of semi-batch recirculation, provided fresh isobutylene of a moderate mol fraction is supplied, there is no equilibrium limit to conversion in water. As TBA is produced the solubility of isobutylene is increased. Obviously it is impractical to operate in such a manner industrially with little conversion of isobutylene (continuously high mol fraction) and batch wise complete recirculation of the liquid phase.

Chapter 7: Synthesis of TBA... 330 Table 7.06 Calculated Measures of the Experimental Design of the CFBR Recirculation Study Run Factors Measures

A B C rTBA η0,TBA rMET η0,MET ξ -1 -1 -1 -1 -1 yIB QL (L.min )xEG mol.g .s mol.g .s 1 0.3 223 0.3 1.817E-08 0.1276 3.926E-09 0.0281 4.6276 2 0.6 223 0.3 3.432E-08 0.1205 7.715E-09 0.0276 4.4483 3 0.3 101 0.3 2.637E-09 0.0185 1.153E-08 0.0826 0.2287 4 0.6 101 0.3 1.489E-08 0.0523 1.568E-08 0.0561 0.9498 5 0.3 223 0.6 2.003E-08 0.2055 8.499E-09 0.0181 2.3563 6 0.6 223 0.6 3.349E-08 0.1718 1.046E-08 0.0111 3.2020 7 0.3 101 0.6 1.466E-08 0.1504 1.200E-08 0.0256 1.2215 8 0.6 101 0.6 8.467E-09 0.0434 1.448E-08 0.0154 0.5847 9(cp) 0.45 162 0.45 9.003E-09 0.0465 7.984E-09 0.0219 1.1277 10(cp) 0.45 162 0.45 8.257E-09 0.0426 7.603E-09 0.0209 1.0860 11(cp) 0.45 162 0.45 1.459E-08 0.0754 6.463E-09 0.0177 2.2580 median 1.653E-08 0.1241 1.100E-08 0.0266 1.7889 avg 1.833E-08 0.1113 1.054E-08 0.0331 2.2023 avg with cp 1.623E-08 0.0959 9.668E-09 0.0296 2.0082

Chapter 7: Synthesis of TBA... 331 7.2-5 Analysis and Discussion of Recirculation Study The Recirculation Study allowed for the comparison of the relative rates of isobutylene reaction with water and ethylene glycol. The ANOVA of the measures of interest of TBA reaction rate, TBA overall effectiveness, MET reaction rate, MET overall effectiveness and selectivity factor are discussed in turn. The effects of the factors are then summarised.

For the observed reaction rate of TBA formation, the ANOVA of Table 7.07 is calculated. For the purpose of this study factors and or interactions returning a Pi of 0.10 or less were considered significant (these values are printed in bold). Thus, with an α value of 10 % the confidence level chosen was that of 90 %. The significant variables were found to be B liquid flowrate, A isobutylene mol fraction. Liquid flowrate was dominant showing the strong response of the reaction rate to changes in the magnitude of external mass transfer rates. As indicated by the probability of SSpq, the extent of curvature was important, indicating a high degree of non-linearity across the range of the study. The observed rate of TBA and MET production were of the same order of magnitude generally.

Table 7.07 TBA Reaction Rate, rTBA, ANOVA VAR Effect Contrast SS DOF MS F P A 8.92E-09 3.57E-08 1.59E-16 1 1.59E-16 13.270 0.068 B 1.63E-08 6.53E-08 5.34E-16 1 5.34E-16 44.514 0.022 C 1.66E-09 6.63E-09 5.49E-18 1 5.49E-18 0.458 0.568 AB 5.89E-09 2.35E-08 6.93E-17 1 6.93E-17 5.781 0.138 AC -5.28E-09 -2.11E-08 5.58E-17 1 5.58E-17 4.657 0.164 BC -1.14E-09 -4.56E-09 2.60E-18 1 2.60E-18 0.217 0.687 ABC 3.94E-09 1.58E-08 3.10E-17 1 3.10E-17 2.589 0.249 SS pq 1.30E-16 1 1.30E-16 10.829 0.081 SS e 2.40E-17 2 1.20E-17 SS total 1.01E-15 10 (Note: A - yIB, B – QL, C - xEG)

The TBA overall effectiveness factor was influenced most dramatically by flowrate. It was also influenced both significantly and positively by the concentration of the solvent ethylene glycol. Ellenberg and Krishna (1999) showed that the separation efficiency of structured packing, which also relies heavily on film type flow, diminishes with increasing surface tension of the liquid due to decreased wetting. The surface tension of

Chapter 7: Synthesis of TBA... 332 ethylene glycol is 44×10-03 N.m compared to 65×10-03 N.m for water, at 348 K. The improved wetting, due to lowered surface tension, works together with higher flowrates to improve the transport of isobutylene. This finding is consistent with the results of the single pass study, which suggests that at higher ReL internal mass transfer can become dominant and external mass transfer limitation can be eliminated. However, the interaction of A isobutylene mole fraction and C ethylene glycol mole fraction, AC, was negative. Although ethylene glycol improves isobutylene solubility and mass transfer, its dilution of water and competitive reaction with isobutylene have a negative effect upon the ability of TBA observed reaction rate to approach intrinsic rates. The solvent thus has numerous opposing effects, which may lead to a very narrow range of feasible operation in any final CED application.

Table 7.08 TBA Overall Effectiveness Factor, η0, ANOVA VAR Effect Contrast SS DOF MS F P A -0.028 -0.114 0.002 1 1.62E-03 5.076 0.1530 B 0.090 0.361 0.016 1 1.63E-02 50.872 0.0191 C 0.063 0.252 0.008 1 7.95E-03 24.865 0.0379 AB 0.008 0.032 0.000 1 1.32E-04 0.411 0.5870 AC -0.042 -0.167 0.004 1 3.50E-03 10.948 0.0805 BC 0.002 0.006 0.000 1 4.75E-06 0.015 0.9141 ABC 0.029 0.114 0.002 1 1.63E-03 5.094 0.1526 SS pq 0.007 1 6.95E-03 21.727 0.0431 SS e 0.001 2 3.20E-04 SS total 0.039 10 (Note: A - yIB, B – QL, C - xEG)

The rate of conversion of the batch of liquid during the experiments was calculated using the batch design equation in the form:

dX Nrm=− (7.22) AA0 dt

The reaction rate and volume were assumed to remain constant during reaction. The mean residence time was calculated from the RTD data for bale packing of Chapter 5

(based on H0L.Bale Equation (5.59)) giving:

−0.7363 tmLB,, =1639.1ReL (7.23)

Chapter 7: Synthesis of TBA... 333 As evident from Table 7.09 the per-pass conversion in the components water and -06 ethylene glycol was very small ie. XW.avg = 4.74×10 .

Table 7.09 Rate of conversion and an estimate of conversion per pass

Run ReL tm dXW/dt XW/pass dXEG/dt XEG/pass s s-1 s-1 1 79 66 3.11E-08 2.04E-06 1.46E-08 9.55E-07 2 79 66 5.87E-08 3.85E-06 2.87E-08 1.88E-06 3 36 118 4.52E-09 5.31E-07 4.28E-08 5.04E-06 4 36 118 2.55E-08 3.00E-06 5.82E-08 6.85E-06 5 47 96 7.66E-08 7.36E-06 2.68E-08 2.57E-06 6 47 96 1.28E-07 1.23E-05 3.30E-08 3.17E-06 7 21 173 5.61E-08 9.67E-06 3.78E-08 6.53E-06 8 21 173 3.24E-08 5.59E-06 4.57E-08 7.88E-06 9(cp) 43 103 2.13E-08 2.19E-06 2.72E-08 2.81E-06 10(cp) 43 103 1.95E-08 2.01E-06 2.59E-08 2.67E-06 11(cp) 43 103 3.45E-08 3.56E-06 2.20E-08 2.27E-06 avg 45 110 4.44E-08 4.74E-06 3.30E-08 3.88E-06

The average production rate for MET was 1.054×10-08 mol.g-1.s-1, which is just below that of the TBA formation rate of 1.833×10-08 mol.g-1.s-1. However, as may be seen from Table 7.10, increasing liquid flowrate has a negative effect on MET formation rate. This is probably due to the shorter residence time in the liquid phase suggesting that the precursor involved in the formation of MET required longer surface relaxation time.

Interestingly for the reaction rate of MET, B liquid flowrate has a negative effect, Table 7.10. This suggests that the rate is controlled by an internal mass transfer step over the range of the study.

The positive interactions BC may stem from the fact that higher amounts of ethylene glycol increased isobutylene solubility, but only through better mixing and liquid renewal is the isobutylene is delivered to the reaction zone on the other side of the fibreglass bags. However, as TBA reaction rate is not affected in the same manner it is possible that MET reaction is affected by a strong product inhibition effect which could be of the form:

Chapter 7: Synthesis of TBA... 334 • Slow MET internal mass transfer and strong adsorption retain MET raising its local concentration and promoting the reverse reaction • The inhibition effect of TBA widely reported for isobutylene hydration also slows down the etherification reaction. Product inhibition is a common trait of isobutylene hydration and etherification, [Velo et al. (1988b) and Parra et al. (1995)].

The diffusion of MET from the site of reaction to the exterior of the catalyst pellet may be slow due to the presence of TBA, and high, local concentrations of MET may have an inhibition effect. Velo et al. (1990) found that for isobutylene, intraparticle diffusivity was found to increase with temperature and decrease with TBA concentration. Values of internal transport were reported to have changed by a factor of 3.5 in a TBA rich environment from those within a pure water. It is possible that TBA and MET affect each other’s intra-particle transport rates in a complex manner.

For the concentrations of the study MET may have begun to react with isobutylene to a significant extent forming DET. The concentration did not exceed 0.008 mol.L-1 and could be considered low. Moreover, significant new peaks were not found in the range of retention times which were expected of bi-tert-butyl ethylene glycol ether (BET). It is possible that MET remains highly associated with the resin, thus producing a high local product concentration and reverse reaction.

Table 7.10 MET Reaction Rate ANOVA VAR Effect Contrast SS DOF MS F P A 3.09E-09 1.24E-08 1.92E-17 1 1.92E-17 30.578 0.0312 B -5.77E-09 -2.31E-08 6.67E-17 1 6.67E-17 106.485 0.0093 C 1.65E-09 6.58E-09 5.42E-18 1 5.42E-18 8.654 0.0987 AB -2.20E-10 -8.81E-10 9.69E-20 1 9.69E-20 0.155 0.7320 AC -8.74E-10 -3.49E-09 1.53E-18 1 1.53E-18 2.438 0.2588 BC 2.01E-09 8.05E-09 8.10E-18 1 8.10E-18 12.935 0.0694 ABC -4.11E-11 -1.64E-10 3.37E-21 1 3.37E-21 0.005 0.9482 SS pq 2.22E-17 1 2.22E-17 35.383 0.0271 SS e 1.25E-18 2 6.26E-19 SS total 1.24E-16 10 (Note: A - yIB, B – QL, C - xEG)

The MET overall effectiveness was very low at an average of 0.03 for the study (Table 7.11). The result was noisy, giving a significant three way interaction, and showing all factors to be significant yet having a negative effect, which in turn suggests that not one

Chapter 7: Synthesis of TBA... 335 of the variables studied was significant, Table 7.11. This strengthens the likelihood that internal mass transport is responsible and that the true effectiveness of MET decreases sharply with catalyst size.

Table 7.11 MET Overall Effectiveness Factor ANOVA VAR Effect Contrast SS DOF MS F P A -1.10E-02 -4.40E-02 2.42E-04 1 2.42E-04 51.386 0.0189 B -2.37E-02 -9.47E-02 1.12E-03 1 1.12E-03 237.838 0.0042 C -3.11E-02 -1.24E-01 1.93E-03 1 1.93E-03 409.008 0.0024 AB 7.28E-03 2.91E-02 1.06E-04 1 1.06E-04 22.483 0.0417 AC 2.46E-03 9.84E-03 1.21E-05 1 1.21E-05 2.563 0.2505 BC 1.78E-02 7.13E-02 6.35E-04 1 6.35E-04 134.551 0.0074 ABC -5.70E-03 -2.28E-02 6.49E-05 1 6.49E-05 13.757 0.0656 SS pq 3.64E-04 1 3.64E-04 77.133 0.0127 SS e 9.44E-06 2 4.72E-06 SS total 4.49E-03 10 (Note: A - yIB, B – QL, C - xEG)

The selectivity factor had an average of 2, and a maximum of 4.6 occurring at high flowrates and low ethylene glycol. Consistent with previously discussed measures selectivity increased with flowrate and decreased with ethylene glycol concentration. However, the ANOVA analysis only suggested liquid flowrate to be significant, Table 7.12.

Table 7.12 Selectivity Factor based upon Observed Reaction Rates ANOVA VAR Effect Contrast SS DOF MS F P A 0.188 0.751 7.04E-02 1 7.04E-02 0.159 0.7284 B 2.912 11.649 1.70E+01 1 1.70E+01 38.371 0.0251 C -0.722 -2.890 1.04E+00 1 1.04E+00 2.361 0.2642 AB 0.146 0.582 4.24E-02 1 4.24E-02 0.096 0.7861 AC -0.083 -0.333 1.39E-02 1 1.39E-02 0.031 0.8758 BC -1.036 -4.145 2.15E+00 1 2.15E+00 4.859 0.1583 ABC 0.596 2.383 7.10E-01 1 7.10E-01 1.606 0.3327 SS pq 1.11E+00 1 1.11E+00 2.500 0.2546 SS e 8.84E-01 2 4.42E-01 SS total 2.30E+01 10 (Note: A - yIB, B – QL, C - xEG)

The selectivity factor was generally higher than that predicted by Equation (4.67) based upon rotating basket reactor kinetics, as demonstrated by Figure 7.08. Runs 1 and 2 exhibited the highest improvements in selectivity and were typified by low ethylene glycol molar fraction (0.3) and high flowrate. Runs 6 to 8 of higher ethylene glycol

Chapter 7: Synthesis of TBA... 336 molar fraction naturally showed typically lower selectivity, however, still surpassed those of the kinetic study. The vertical spread in selectivity was due mainly to variation in flowrate. For the range of this study isobutylene concentration was not a significant variable in terms of selectivity factor. As isobutylene is unlikely to be adsorbed onto the surface, and each reaction is first order in isobutylene the lack of dependence of selectivity upon isobutylene concentration follows. The same lack of dependence was observed in the kinetic study.

Chapter 7: Synthesis of TBA... 337 5 01 02 4

06 3 CFBR Study ξ

05 cp 2

07 1 04

Selectivity Factor 08

03 0 012345 Selectivity Factor ξ Kinetic Study

Figure 7.08 A comparison of selectivity factors obtained in the Kinetic Study and CFBR runs, showing the run id of the experimental design.

The effect of higher concentrations of TBA and MET upon isobutylene liquid phase and effective diffusivity have yet to be quantified, thus an analysis of the form used in Section 7.2-3 was not considered adequate. However, using the general form of Equation (7.01), and the solubility of isobutylene, it was possible to calculate the total observed resistance cf. Table 7.13. It was found that Rtot,obs for TBA was a strong function of flowrate, however, the total resistance for MET had almost the opposite behaviour. This again suggested that the effects internal to the catalyst may be more significant and further that the presence of TBA influences these. Figure 7.09 shows the influence of observed TBA reaction rate upon observed MET resistance, Rtotal. TBA reaction rate was significantly influenced by isobutylene molar fraction, yIB, thus the plot was constructed as two series. It was found that indeed the overall resistance to MET reaction rate increased with TBA observed reaction rate. This has some form of inhibiting effect upon MET production and explains the significance of nearly all variables of the ANOVA for MET overall effectiveness, Table 7.11.

Chapter 7: Synthesis of TBA... 338 Table 7.13 Total Observed Resistance for MET and TBA CFBR recirculation runs.

Run Factors Rtot.observed A B C Rtot.obs Rtot.obs -1 yIB QL (L.min )xEG TBA Met 1 0.3 223 0.3 57755 267264 2 0.6 223 0.3 61147 272000 3 0.3 101 0.3 397810 90968 4 0.6 101 0.3 140898 133820 5 0.3 223 0.6 73196 172471 6 0.6 223 0.6 87544 280315 7 0.3 101 0.6 100002 122156 8 0.6 101 0.6 346259 202442 9(cp) 0.45 162 0.45 209510 236275 10(cp) 0.45 162 0.45 228445 248088 11(cp) 0.45 162 0.45 129263 291869

300000

250000

200000

tot.observed 150000

100000 MET R yIB 0.3 50000 yIB 0.6

0 0 1E-08 2E-08 3E-08 4E-08 -1 -1 rTBA (mol.g .s )

Figure 7.09 MET observed overall resistance as a function of observed TBA

reaction rate, rTBA.

Chapter 7: Synthesis of TBA... 339 The following argument was proposed for the responsiveness of TBA reaction rate to flowrate, improved isobutylene mass transfer and the near inverse behaviour of MET reaction rate. Both water and ethylene glycol adsorb onto the active sites, as indicated by the kinetic model Equation (4.76). Water preferentially swells the catalyst and occupies more sites, as indicated by the inhibition of the reaction between ethylene glycol and isobutylene through active site hydration. Ethylene glycol occupies less sites however its ability to attract isobutylene is stronger. Ethylene glycol promotes isobutylene solubility and marginally improves mass transfer coefficients, which are most improved through higher flowrates and better bale liquid renewal as indicated by lower Per. The isobutylene reacts preferentially with the adsorbed ethylene glycol, thus the production of MET is less affected by the rate of external mass transfer and that of hydration more so. The ability of MET to leave the catalyst is poorer than that of TBA. MET is retained by poorer diffusivity and adsorbs more readily. Further, its effective diffusivity may be negatively affected by TBA, as indicated by a trend of its overall resistance in TBA reaction rate. In its adsorbed form it reverts back to ethylene glycol and isobutylene, alternatively it begins to take part in a consecutive reaction, which produces an even heavier species..

Thus, the conditions, which promote TBA selectivity, are:

• Minimal use of entrainer • High external mass transfer coefficients for IB

Further, conditions which may promote selectivity factor, but require further investigation are:

• Larger size of catalyst • Higher TBA concentration. TBA may be detrimental to the effective diffusivity of MET. While tortuosity is a property of the catalyst, the interaction of MET and the resin may be affected by the presence of TBA. Such behaviour has been reported for the case of isobutylene diffusion, Velo et al (1990). • Higher temperatures. For the rotating basket reactor selectivity factor improved with temperature, as the responsiveness of TBA reaction rate was higher thus its

Chapter 7: Synthesis of TBA... 340 activation energy lower. Higher temperatures may shift MET equilibrium to the left to a greater extent than for the case of TBA.

The proposed CED process may easily accommodate some of these features as well as separation effects to further improve the selectivity factor. Minimal use of an entrainer to promote selectivity factor will be a dominant consideration. High reflux rates can be used and a narrow extractive distillation section packed with high efficiency packing in order to minimise the requirement of high entrainer ratio. The high reflux rate will also keep TBA present and potentially diminish the rate of the side reaction.

7.3 Synthesis of TBA by Catalytic Extractive Distillation over Bale Packing Catalytic distillation is an attractive process as it can assist in improved conversion beyond equilibrium limitation. The obstacles to the implementation of catalytic distillation for the synthesis of TBA were identified as:

1. the existence of a particularly pressure insensitive azeotrope between reactant water and product TBA at 353 K, 65 mol % TBA and 1 atm. 2. the lack of solubility of isobutylene in water

A catalytic extractive distillation column design of promotion of natural volatility between water and TBA, using ethylene glycol was proposed as a potential solution, which may lower these hurdles towards the combination of separation and reaction.

The extractive distillation study of Chapter 6 showed that the azeotrope between water and TBA is readily crossed through the use of ethylene glycol as entrainer. The minimum non-pinching solvent to feed molar fraction was predicted as xr = 0.165 and experimentally it was demonstrated that xr = 0.42 was optimum. Reverse entrainment was observed beyond xr ≈ 0.55.

Operating under CFBR conditions it was found that TBA reaction rates improved with solvent through the increased availablity of isobutylene. This improved supply of

Chapter 7: Synthesis of TBA... 341 reactant arose both due to both improved equilibrium solubility and improved mass transfer coefficients from the gas phase to the solid. This improvement was gauged through predicted and measured resistances to reaction. However, comparable amounts of MET were formed under these conditions. It is proposed that the CED design with distillation separation and elevated temperatures would allow for better TBA purity and higher reactive section selectivity factor.

The experiments were conducted at atmospheric pressure with zero reflux utilising a column stream setup closer to absorption than true distillation. This arrangement had been utilised in the extractive distillation study. These runs were extended with a series of runs at total reflux. Under these conditions it was possible to carry out solvent free runs for direct comparison.

7.3-1 Experimental The separate components of the CFBR and Extractive Distillation experiments were brought together for the combination of reaction and separation within one apparatus referred to as Catalytic Extractive Distillation. The experimental rig of Figure 7.10 was developed. The distillation was performed at atmospheric conditions. The parameters manipulated include water boilup rate, solvent feed rate, reflux ratio gas feed composition.

7.3-1-1 CED Procedure The CED for the synthesis of TBA experimental rig was presented in Figure 7.10. The rig employed was similar to that used for extractive distillation experiments, however incorporated provisions for the addition of isobutylene as a dilute gas and for the online analysis of isobutylene. Isobutylene and nitrogen were supplied by rotameter as for CFBR experiments and analysed by GC as described in Section 7.2-1-1. Additionally, a different condenser arrangement was used to that used for extractive distillation experiments. The condenser used had a smaller dead volume ensuring that the small quantities of TBA produced could be collected. The outlet arrangement allowed for the non-condensed gas phase to be collected and sent to a GC split arrangement allowing for a constant low flowrate of gas to be sampled. The top liquid collection vessel (not shown) was maintained at the same pressure as the gas line and its volume calibrated for

Chapter 7: Synthesis of TBA... 342 flow monitoring. The condenser water was maintained at such a level (2 to 3 L.min-1) as to maintain atmospheric conditions determined through the use of a manometer. The distillate produced was subcooled to typically 313K.

The flowrate at the bottom of the column, at the reboiler bypass was monitored hot using a calibrated volume and stopwatch. The flowrate of the entrainer ethylene glycol was measured by hot calibrated rotameter. Column temperature was measured at 6 key points as shown using mineral insulated K-type thermocouples, a high accuracy temperature readout and thermocouple selector switch.

Liquid samples were taken from the distillate and the column bottom streams. These were analysed for TBA, MET, water, DET, and DIB. Samples along the column length were not taken as:

• Bale packing is difficult to sample effectively • The bed of bale packing was relatively short and the conversion low

The column was heated through boilup of steam under conditions of total reflux. Ethylene glycol was next introduced and the bottom bypass adjusted. Once the temperature profile had stabilised the gas side reactant isobutylene, diluted with nitrogen, was introduced at a total flowrate of 0.4 L.min-1. The column was placed in 0 reflux for complete distillate collection and operation in a mode similar to absorption as described in Chapter 6. The column was allowed to stabilise for 45 min after which sampling was commenced and five samples collected at 15 minute intervals.

The reboiler had deliberately been chosen to be large so as to maintain a constant feed rate of steam. The 10 L reboiler was charged with 8 L of distilled water to a pre- calibrated marker at the beginning of each experimental run. The two to three hours of experimental operation resulted in an average volume change of 2 L. The calibration of the reboiler flow rate was checked before and after the experimental runs and boilup rates were found to remain constant.

Chapter 7: Synthesis of TBA... 343 [04]

P [03] P

T T1 T [02] [01]

T [07] T2

T3 [05] [06] T7

T T4

[08][09] [10] T

T8

T5 ETHYLENE GLYCOL TANK T6 SILICON OIL TANK [11] [12]

[13]

BOTTOMS TANK

[14] P P

[15]

Figure 7.10 CED Experimental Rig

Key to Figure 7.10 CED Experimental Rig [01] Cooling Water Rotameter.

Chapter 7: Synthesis of TBA... 344 [02] Condenser, Single Coil Unjacketed, QuickFit 19/26, Crown Scientific. [03] Gas Sampling Assembly, 6 Way Valco Valve, Whitey Needle Valves, Soap bubble meter 100 mL. [04] GC-8A TCD, Shimadzu. [05] Temperature Indicator and Thermocouple Selector Switch, Shinko Controller, RS Selector. [06] Column: Jacketed 45 mm id, 60 mm od. [07] GC Integrator. [08] Liquid Feed Preheater: Silicon oil bath with Braun Bath Heater, SS bath, Heat Exchanger Coil, ¼” SS 40 Rectangular Loops at 40 cm/loop. [09] Liquid Feed Rotameter: Planton, Duff and Machintosh, Model 48, universal, with scale marked for 2 to 25 L/min Air. [10] Jacket Heating Oil Bath: Bath Heater and Recirculator, Haake DC 5; Insulated SS tank. [11] Liquid Feed Pump: Masterflex Peristaltic Pump, Variable Speed Variable Speed Modular Drive, Model 7016 pump head, L/S 16 Viton Tubing, Cole Parmer. [12] Bottoms Liquid Split with calibrated liquid volume. [13] Glass Reboiler, 10 L or 5 L and Heating Mantle, 10 L. [14] Mantle Power regulator. [15] Gas Supply: Built inhouse in a stable/sturdy aluminium frame. Includes: 3 × 10” Rotameters: Fisher Porter Meter Tube No FP ¼-20-G-5/84 Meter Tube No FP 1/8-12-G-5/84 Meter Tube No FP 1/8-08-G-5/84 1 × 6” Rotameter: Fisher Porter Meter Tube No FP 1/8-08-P-3/37 2 × Gas Regulator: Fairchild Regulator, Kendall, Model 10, Range 5 to 400 PSI 2 × Pressure Gauges: Dobbie Instrument Australia 0 to 250 kPa 2 × Needle Valves, Whitey SS ¼”

Chapter 7: Synthesis of TBA... 345 Stepwise CED Experimental Procedure The following procedure was used for the CED experiments: 1. Prepare GC-8A FID, (need at least 2 hr for this step). 2. Prepare GC-8A TCD. 3. The column should be packed with catalytic bales containing Amberlyst 15 in the 500 to 700 μm particle size range. Other parts of the column to be packed with 8 mm raschig rings. 4. The cooling water [1] is opened and adjusted to 2 to 3 L/min, giving atmospheric conditions checked manometrically. 5. Charged the reboiler [13] with distilled water and bring to boil by adjusting the mantle power supply [14]. 6. Prepare EG Tank Charge, make sure there is sufficient room in the waste container, prepare a spare if necessary. 7. Turn on the Silicon Oil Recirculator and set to the desired value, process + 15 oC. 8. Turn on the EG pump [11] and set the flow rate to that required using the rotameter [09], turn the three way valve towards the column to begin the experimental run. The solvent flowrate will need to be adjusted at this point to compensate for the higher-pressure drop of pumping into the column. 9. Monitor and open the bottoms bypass [12]. Turn off the flow and measure the bottoms flow rate using the calibrated volume and a stopwatch. 10. Adjust the feed gas Rotameters [15] to desired values using predefined calibrations. 11. Commence timing of the experiment once the isobutylene is introduced. 12. Take the temperature profile at regular intervals using the selector switch and temperature indicator [05]. Watch this profile by plotting for the approach to steady state. 13. Take top and bottom samples (5) every 15 minutes after an initial 45 min stabilisation period. 14. To end run turn off the EG, top up the reboiler water and turn to max boil up rate and wash down the column. Set the new setpoints of the next run. 15. For complete shutdown turn off reboiler and preheaters. Allow the column to cool down and then turn off cooling water

Chapter 7: Synthesis of TBA... 346 7.3-1-2 Experimental Design As in the case of CFBR runs a full factorial centre point experimental design was chosen for the bulk of experimentation conducted at zero reflux. The variables studied included water boilup rate, solvent feed rate and isobutylene feed molar fraction, Table 7.14. These were considered the key variables of the process, subsequent to extractive distillation and CFBR experimentation.

Table 7.14 Experimental Design Factor Levels of CED Study Main Runs Factor Definition Lower Higher Level Units Level (-1) (+1) -1 A Water Boilup Rate, F1 0.30 0.60 mol.min -1 B Solvent Feed Rate, F3 0.21 0.65 mol.min -1 C IB Feed Fraction, yIB 0.2 0.44 mol.mol

The range of the boilup rate and solvent feed rate was based upon the results of the extractive distillation study of Chapter 6. They ensured that the column was well wetted operated across a wide range within the pre-loading regime and that relevant solvent to feed ratios were explored.

Table 7.15 Solvent to water ratios obtained across the study Run W S S/W mol.min-1 mol.min-1 1 0.30 0.21 0.71 2 0.60 0.21 0.35 3 0.30 0.65 2.17 4 0.60 0.65 1.07 5 0.30 0.21 0.71 6 0.60 0.21 0.35 7 0.30 0.65 2.17 8 0.60 0.65 1.07 cp 0.45 0.41 0.90

These main runs were extended with additional runs conducted at total reflux investigating maximum column potential and MET formation. These additional runs allowed a comparison to solvent free operation at least in terms of product purity. Solvent-free operation was not possible under conditions of zero reflux as no liquid return would be provided and the column would not function properly.

Chapter 7: Synthesis of TBA... 347 7.3-2 Results and Discussion The results obtained can be referenced to the stream diagram of Figure 7.11. The compositions and flow of the inlet streams 1, 2 and 3 were known experimental setpoints. However, the gas feed stream 2 was additionally analysed and these values used in calculations. The outlet stream (4, 5 and 6) composition was unknown and determined by GC analysis and material balance.

TBA was detected in the distillate (stream 5) and the levels in general were reasonable given the short bed and continuous mode of operation. The average concentration attained in the top product stream was 7.4 mmol.L-1, with a range of 1.8 to 16.9 mmol.L-1 over the main study. Most significantly, no TBA was detected in the bottom stream 6. This was very encouraging as it suggested good product separation and little product entrainment in the downward moving liquid phase.

4. Non-Condensables IB, N2

Condenser 5. Distillate / Product W, TBA, MET, IB

3. Ethylene Glycol

CED Column

Bale Packing

2. Gas Feed IB, N2

1. Steam 6. Bottoms EG, W, MET, IB

Figure 7.11 CED Column Stream Number Specification

Chapter 7: Synthesis of TBA... 348 Table 7.16 Summary of Dependant and Independent CED Study Variables, Main (1 to 11) and Additional (12 to 17) Runs. For Stream Identification refer to Figure 7.11. Variable A C B Stream 1 2 3 4 5 6 Exp Reflux F1.W yIB F3.EG yIB Q5 CTBA CMET Q6 CW mol/min mol/min L/min mol.L-1 mol.L-1 L/min mol.L-1 Method: setpoint setpoint setpoint/ setpoint GC-TCD-G measured GC-FID-L GC-FID-L measured GC-TCD-L measured 1 0 0.298 0.144 0.200 0.1427 0.0032 0.0053 0.0 0.0041 48.7221 2 0 0.601 0.139 0.200 0.1336 0.0098 0.0018 0.0 0.0044 52.8981 3 0 0.298 0.145 0.614 0.1397 0.0032 0.0058 0.0 0.0111 24.7444 4 0 0.601 0.136 0.614 0.1187 0.0084 0.0041 0.0 0.0088 29.1850 5 0 0.298 0.431 0.200 0.4242 0.0039 0.0107 0.0 0.0051 45.2031 6 0 0.601 0.429 0.200 0.4112 0.0080 0.0075 0.0 0.0055 47.4352 7 0 0.298 0.436 0.614 0.4150 0.0030 0.0169 0.0 0.0079 35.6465 8 0 0.601 0.436 0.614 0.4097 0.0083 0.0078 0.0 0.0098 26.0660 9(cp) 0 0.449 0.299 0.385 0.2959 0.0061 0.0055 0.0 0.0062 37.4851 10(cp) 0 0.449 0.292 0.385 0.2455 0.0060 0.0065 0.0 0.0067 43.1369 11(cp) 0 0.449 0.292 0.385 0.2821 0.0053 0.0061 0.0 0.0067 39.0201 12 ∞ 0.298 0.027 0.200 0.0286 - 0.1996 0.0 0.0083 40.7426 13 ∞ 0.601 0.026 0.200 0.0247 - 0.0527 0.0002 0.0187 45.4653 14 ∞ 0.298 0.014 0.614 0.0109 - 0.2327 0.0001 0.0169 22.1288 15 ∞ 0.601 0.026 0.614 0.0255 - 0.0710 0.0 0.0250 25.0273 16 ∞ 0.298 0.018 0.000 0.0180 - 0.0069 0.0 0.0065 - 17 ∞ 0.601 0.019 0.000 0.0183 - 0.0116 0.0 0.0134 -

Chapter 7: Synthesis of TBA... 349 Table 7.17 CED Column Temperature Profiles RUN T1 T2 T3 T4 T5 T3to5 avg T6 K K K K K K K 1 374 376 374 377 372 374 378 2 377 375 379 372 370 374 377 3 371 373 375 382 372 376 377 4 373 383 373 373 372 373 378 5 375 373 373 375 372 373 378 6 379 376 374 378 373 375 377 7 373 373 383 379 364 376 376 8 380 381 378 378 368 375 379 9(cp) 377 375 383 378 374 379 377 10(cp) 375 384 386 377 370 378 376 11(cp) 375 378 385 376 369 377 376 12 372 370 374 373 368 372 377 13 374 372 376 376 373 375 380 14 372 372 378 380 368 375 379 15 374 373 378 377 362 372 379 16 372 370 372 371 368 370 376 17 373 372 373 373 370 372 378

Chapter 7: Synthesis of TBA... 350 The following material balances were written for the components of the CED study. The discrepancy, β, could be estimated for those components whose flows and concentrations were completely specified.

Steady-state component balances in component molar flow Fst.i, were proposed with stated assumptions. TBA was not observed in stream 4 or 6 over the entire study F4.TBA

= F6.TBA = 0, thus:

TBA: m.rTBA = F5.TBA (7.24)

MET was not detected in any of the outlet streams. If MET was formed it remained below its limit of detection of:

• 0.002.0 mmol.L-1 analysis of solvent rich stream 6 • 0.1 mmol.L-1 analysis of solvent free distillate stream 5

The potential rate of MET formation was thus estimated based upon the isobutylene balance.

For the purpose of the isobutylene balance it was assumed that the liquid outlet streams were saturated with isobutylene. The temperature of stream 5 was 313 K.

IB: F2.IB = F4.IB + F5.IB +F6.IB + m.rTBA + m.rMET + βIB (7.25)

The balance residual, β, and actual rate of MET formation could be decoupled and their sum was used as the estimate of possible rMET. Back calculation of the concentration of MET expected in stream 6 showed that this estimate was higher than allowed by the limit of MET detection. The estimate was used as an indicator that MET reaction is possible. Other contributions to the overall value of βIB may include loss, unaccounted for side reaction and analysis error.

Chapter 7: Synthesis of TBA... 351 Ethylene glycol was preheated and checked for water content, which was found to be insignificant, thus the solvent feed stream was considered water free, F3.W = 0. A balance over water gave:

W: F1.W = F6.W + F5.W + m.rTBA + βW (7.26)

Table 7.19 suggested that the GC calibration overestimated the content of water of stream 6 and a balance residual was observed. Further calculations were based upon measured flowrates.

Ethylene glycol was not detected in the top streams, thus there was no carry over and the following balance could be written:

EG: F3..EG = F6.EG (7.27)

The observed reaction rate of TBA as reported in Table 7.18 was analysed according to the experimental design and the ANOVA reported Table 7.20. As may be expected, each of the investigated variables was significant. More interestingly over the range of study there was no interaction and near linearity was achieved, each of the factors had a positive effect. Isobutylene molar fraction was found to be the most dominant among the factors. Any improvement in its transport and solubility provides for substantial improvement, which truly reflects its linear relationship with the reaction rate and its limiting nature within the reaction network. The effect of ethylene glycol and the positive effects of all factors are depicted in Figure 7.12. Given that there were no interactions and that the degree of curvature across the study range as quantified by SSpq was not significant, over the range of the study it was possible to obtain a simple linear regression for rTBA.obs:

−−09 09 −09 −09 rFTBA. obs =−1.75 × 10 + 3.1 × 10W + 1.97 × 10FEG + 9.66 × 10 yIB (7.28)

Chapter 7: Synthesis of TBA... 352 Table 7.18 Isobutylene Balance on Main Runs under Zero Reflux

Exp F2IB F4IB F5 IB F6 IB m.rTBA (β+m.rMET) rTBAobs rMET CMET.6 mol.min-1 mol/min mol/min mol/min mol/min mol/min mol/g.s mol/g.s mol.L-1 1 2.36E-03 2.33E-03 1.21E-06 2.45E-06 1.63E-05 7.83E-06 1.47E-09 7.06E-10 0.0019 2 2.27E-03 2.18E-03 3.47E-06 2.93E-06 1.73E-05 6.29E-05 1.56E-09 5.67E-09 0.0142 3 2.37E-03 2.28E-03 1.18E-06 8.05E-06 1.85E-05 5.62E-05 1.67E-09 5.07E-09 0.0051 4 2.22E-03 1.94E-03 2.64E-06 5.89E-06 3.46E-05 2.35E-04 3.12E-09 2.11E-08 0.0267 5 7.05E-03 6.93E-03 4.42E-06 9.91E-06 4.19E-05 5.95E-05 3.78E-09 5.36E-09 0.0116 6 7.01E-03 6.72E-03 8.72E-06 9.03E-06 5.57E-05 2.17E-04 5.02E-09 1.96E-08 0.0397 7 7.12E-03 6.78E-03 3.34E-06 1.72E-05 5.12E-05 2.64E-04 4.61E-09 2.38E-08 0.0334 8 7.12E-03 6.70E-03 9.01E-06 2.09E-05 6.48E-05 3.29E-04 5.84E-09 2.96E-08 0.0337 09(cp) 4.77E-03 4.61E-03 4.81E-06 8.82E-06 3.39E-05 1.09E-04 3.05E-09 9.82E-09 0.0176 10(cp) 4.77E-03 4.61E-03 3.92E-06 9.22E-06 3.87E-05 1.05E-04 3.49E-09 9.42E-09 0.0157 11(cp) 4.77E-03 4.61E-03 3.99E-06 8.79E-06 3.23E-05 1.11E-04 2.91E-09 1.00E-08 0.0166

Table 7.19 Water Balance on Main Runs under Zero Reflux

Exp F1 W F6 W F5 W F6+F5 βW W Split mol.min-1 mol.min-1 mol.min-1 mol.min-1 mol.min-1 F5/F1 1 0.2977 0.2019 0.1763 0.3781 -0.0804 0.5921 2 0.6011 0.2352 0.5422 0.7773 -0.1762 0.9020 3 0.2977 0.2811 0.1758 0.4569 -0.1592 0.5905 4 0.6011 0.2610 0.4644 0.7253 -0.1242 0.7726 5 0.2977 0.2322 0.2172 0.4494 -0.1517 0.7298 6 0.6011 0.2610 0.4420 0.7030 -0.1019 0.7353 7 0.2977 0.2858 0.1678 0.4536 -0.1559 0.5636 8 0.6011 0.2601 0.4586 0.7187 -0.1176 0.7629 09(cp) 0.4494 0.2339 0.3388 0.5726 -0.1232 0.7538 10(cp) 0.4494 0.2902 0.3327 0.6229 -0.1735 0.7403 11(cp) 0.4494 0.2637 0.2948 0.5585 -0.1091 0.6560

Chapter 7: Synthesis of TBA... 353 -1 The optimum reaction rate over the study was clearly obtained at Fw = 0.6 mol.min , -1 FEG = 0.61 mol.min and yIB = 0.44. The distillate product stream (stream 5) was free of MET. The limit of detection of MET in the absence of solvent is equal to that of TBA of 0.1 mmol.L-1, which lends confidence to its absence. The bottom stream (stream 6) was free of TBA thus extractive distillation effect of enhanced volatility serves to enrich the vapour phase with TBA. This atmospheric investigation suggests that a pressurised investigation with increased isobutylene solubility is warranted.

Table 7.20 TBA Reaction Rate, rTBA.obs, ANOVA Effect Contrast SS DOF MS F P A 1.00E-09 4.01E-09 2.01E-18 1 2.01E-18 22.16 0.0423 B 8.52E-10 3.41E-09 1.45E-18 1 1.45E-18 15.96 0.0573 C 2.86E-09 1.14E-08 1.63E-17 1 1.63E-17 179.69 0.0055 AB 3.35E-10 1.34E-09 2.24E-19 1 2.24E-19 2.46 0.2571 AC 2.32E-10 9.29E-10 1.08E-19 1 1.08E-19 1.19 0.3897 BC -2.80E-11 -1.12E-10 1.57E-21 1 1.57E-21 0.02 0.9076 ABC -3.43E-10 -1.37E-09 2.35E-19 1 2.35E-19 2.58 0.2493 SS pq 1.18E-19 1 1.18E-19 1.30 0.3729 SS e 1.82E-19 2 9.09E-20 SS total 2.07E-17 10 (Note: A – Boilup Rate F1, B – Solvent Feed Rate F3, C yIB.2)

Intrinsic reaction rates were estimated for the same concentration and temperature. The bale reactive section average values of concentration and temperature were used. These were estimated assuming that the separation efficiencies of bale and raschig packing are similar. The solubility of isobutylene was estimated assuming ideal vapour phase behaviour for the dilution of stream 2 containing isobutylene nitrogen by stream 1 composed entirely of steam.

Table 7.21 Intrinsic reaction rate for same temperature and concentrations as the average of the bale reactive zone

Run CW.avg CEG. avg CIB rTBA.int rTBA.obs η0 (Eqn 7.11) mol/g.s mol/g.s 1 6.26 15.05 1.27E-05 1.10E-09 1.47E-09 1.34 2 3.27 16.01 6.16E-06 2.67E-10 1.56E-09 5.85 3 2.25 16.33 1.48E-05 4.35E-10 1.67E-09 3.83 4 2.51 16.25 6.47E-06 2.13E-10 3.12E-09 14.62 5 4.35 15.66 3.84E-05 2.25E-09 3.78E-09 1.68 6 7.87 14.53 1.71E-05 1.90E-09 5.02E-09 2.65

Chapter 7: Synthesis of TBA... 354 7 2.39 16.29 4.47E-05 1.40E-09 4.61E-09 3.29 8 2.61 16.22 2.08E-05 7.13E-10 5.84E-09 8.19 9(cp) 3.19 16.03 1.86E-05 7.84E-10 3.05E-09 3.89 10(cp) 3.35 15.98 1.81E-05 8.04E-10 3.49E-09 4.34 11(cp) 4.33 15.66 1.80E-05 1.05E-09 2.91E-09 2.77

The estimated intrinsic rates suggested an enhancement effect as the overall effectiveness factor was generally greater than unity. Possible reasons for this enhancement included:

• The removal of TBA into the vapour phase, out of the reaction zone promotes reaction rate. This type of enhancement is desirable in the application of catalytic distillation processes. • The intrinsic reaction rate at 373 K is greater than that predicted by the model of Equation (4.76). The rate predicted and that observed were of similar magnitude and were of the lower range of the magnitude observed within the kinetic study, thus very sensitive to any error inherent to the model. • The solubility of isobutylene is higher than estimated, possibly due to the non-ideal effect of steam, which drives isobutylene to dissolve preferentially in the ethylene glycol rich liquid phase. The influence of steam upon solubility of VLE has yet to be determined. The literature review of thermodynamic property sets of Chapter 6 found that isobutylene thermodynamic model parameters associated with this system for any other model than UNIFAC are as yet unavailable. Upon pressurisation it is likely that both the vapour and liquid phases will behave non- ideally.

Chapter 7: Synthesis of TBA... 355 7.E-09

6.E-09 FW 0.3, yIB 0.14

) FW 0.6, yIB 0.14 -1

.s 5.E-09 FW 0.3, yIB 0.44 -1 FW 0.6, yIB 0.44 4.E-09 (mol.g 3.E-09

2.E-09 TBA.observed r 1.E-09

0.E+00 0 0.2 0.4 0.6 0.8 -1 FEG (mol.min )

Figure 7.12 The effect of ethylene glycol feed rate upon observed TBA reaction rate under CED conditions and zero reflux

0.30 FW 0.3 mol/min 0.25 FW 0.6 mol/min

) 0.20 -1

0.15 (mol.L TBA

C 0.10

0.05

0.00 0 0.2 0.4 0.6 0.8 -1 FEG (mol.min )

Figure 7.13 The effect of ethylene glycol feed rate upon observed TBA concentration under CED conditions and total reflux

Chapter 7: Synthesis of TBA... 356 Additional runs were conducted at total reflux and yIB = 0.44, allowing for the comparison of pure water runs to those containing the solvent ethylene glycol. Under conditions of total reflux the effectiveness of ethylene glycol as extractive distillation entrainer may suffer, however they do allow for a base comparison. Trace amounts of MET were detected in the distillate for two of the runs. MET has a boiling point of 425 K, it may have some azeotropic behaviour within the system, which can be affected by the solvent. If it forms an azeotrope with water, it is likely to form a minimum boiling azeotrope of temperature just below 373 K, typical of middle range boiling alcohols and ethers. For the purpose of further development of this process, the physical and thermodynamic properties of MET and DET need to be evaluated, as these are not available in open literature. The purity of the reflux obtained was shown as a function of

FEG and FW in Figure 7.13. Based on Figure 7.13 it is evident that ethylene glycol enhances TBA formation rate and enriches the product stream. Concentrations of TBA, 34 times greater than for pure water runs, could be achieved over the range of this study.

Jimenez and Costa-Lopez (2002a) modelled a reactive extractive distillation column for the transesterification of methyl acetate to butyl acetate over Amberlyst 15 using the structured packing KATAPAK-S. They used the solvent o-xylene and worked with methyl acetate feed to entrainer ratios of 1.8 to 2.2 and reactant ratio of 1:1. For their system miscibility, mass transfer limitations and side reaction were not significant issues or were simply not addressed. As a result their rate based and equilibrium based simulations returned very similar results. In terms of solvent use, they were primarily concerned with crossing the methanol-methyl acetate and butanol and butyl acetate azeotropes. As with the TBA system, these occur between reactant and product thus are not crossed through reaction. The entrainer of this study ethylene glycol has a much greater role thus separation and reaction are coupled to a greater extent. This may lead to the system being over constrained or conflicting requirements being placed upon feed to entrainer ratio. This is especially relevant as side reaction is of concern. Increased ethylene glycol improves isobutylene solubility and mass transfer coefficients, however lowers the selectivity factor. In order to relax the benefits required of the entrainer it may be possible to:

• Improve external mass transfer through better hardware design

Chapter 7: Synthesis of TBA... 357 • Pressurise the system condensing isobutylene, but keep the two resulting phases well mixed in a similar manner as the kinetic studies of Gupta and Douglas (1967) and Ihm et al. (1988a), where these were sheared to the point of emulsion type behaviour.

7.4 Concluding Remarks Under CFBR conditions and recirculation type experiments both MET and TBA were formed in appreciable amounts. The reaction rates observed were similar and limited by the external mass transfer of isobutylene. The mass transfer was improved by higher flowrates in the liquid phase and the use of ethylene glycol.

Under CED conditions, TBA was enriched and generally separated from any MET, which may have been formed. The entrainer, ethylene glycol, had a positive effect upon reaction rate under conditions of zero reflux and absorption type operation. Ethylene glycol improved the availablity of isobutylene and enriched the product stream. Under conditions of total reflux a clear improvement in TBA purity was observed. However, typically, the reaction rates over a short bed of five catalytic bales were inadequate to obtain azeotropic conditions. The direction of process pressurisation should be pursued for improved isobutylene concentration, which remains the limiting factor of the process.

The solvent and entrainer ethylene glycol warrants further investigation, as the product TBA should be able to be isolated using CED technology. The process byproduct MET remains a concern, but its formation seems to be limited under CED conditions.

Hardware for catalytic distillation based upon the rotating basket reactor of Chapter 4 was proposed and its feasibility investigated in the next chapter.

7.5 Nomenclature

2 -3 ae effective area of packing, (m .m ) 2 -3 at effective area (m .m ) 2 -3 ap packing surface area, (m .m )

Chapter 7: Synthesis of TBA... 358 A,B,C Factors of experimental design -1 CAi Concentration of component A in stream i, (mol.L ) dp characteristic diameter, (m) D diffusivity, (m2.s-1) 2 -1 De effective diffusivity, (m .s ) -1 -1 Fi.A molar flowrate of component A in stream i, (mol.min ) or (mol.s ) 3 -3 HL Liquid Phase Total Holdup, volume fraction of reactor volume, (m .m ) 3 -3 HD,L Liquid Phase Dynamic Holdup, (m .m ) 3 -3 HS,L Liquid Phase Static Holdup, (m .m ) k’ pseudo first order reaction rate constant, kGai bulk gas to interface mass transfer coefficient kLai interface to bulk liquid mass transfer coefficient ksap bulk liquid to solid mass transfer coefficient m mass of catalyst, (g) -1 Qi volumetric flowrate of stream i, (L.min ) -1 -1 rA reaction rate of component A, (mol.g .s ) R reaction rate resistance, (g.s.L-1)

Rei Reynolds Number of phase i, = ρiUidr/μi

Reai Reynolds Number Bale packing, = 4Uiρi/atμi

Sci Schmidt Number of phase i

Shi Sherwood Number of phase i U Superficial velocity based on reactor area, (m.s-1) V reactor or liquid volume, (m3) xr solvent to azeotropic feed molar fraction Equation (6.20) xA liquid phase molar fraction of component A yi Experimental Design measure yA gas phase molar fraction of component A Greek: ε bed porosity η effectiveness

η0 overall effectiveness μ viscosity (N.s.m-2) ν experimental design level –1, +1

Chapter 7: Synthesis of TBA... 359 ρ density (kg.m-3) ξ selectivity factor

Subscripts: G Gas L Liquid m mean

Acronyms: CED Catalytic Extractive Distillation CFBR Countercurrent Fixed Bed Reactor dof degree of freedom DET ethylene glycol di-tert-butyl ether EG Ethylene Glycol IB Isobutylene MET ethylene glycol mono-tert-butyl ether MS Mean Square RTD Residence Time Distribution SS Sum of Squares TBA tert-Butyl Alcohol

7.6 Literature Cited Ellenberger, J. and R. Krishna “Countercurrent operation of structured catalytically packed distillation columns: pressure drop, holdup and mixing.” Chem. Eng. Sci. 54: 1339. (1999). Gupta, V. P. and J. M. Douglas “Diffusion and Chemical Reaction in Isobutylene Hydration within Cation Exchange Resin.” AIChE Journal 13(5): 883-889. (1967). Herskowitz, M., R. G. Carbonell and J. M. Smith “Effectiveness Factor and Mass Transfer in Trickle-Bed Reactors.” AIChE Journal 25(2): 272-282. (1979). Hines, A. L. and R. N. Maddox "Mass Transfer, Fundamentals and Applications". New Jersey, P T R Rrentice Hall. (1985).

Chapter 7: Synthesis of TBA... 360 Ihm, S. K., M. J. Chung and K. Y. Park “Activity Difference between the Internal and External Groups of Macroreticular Ion Exchange Resin Catalysts in Isobutylene Hydration.” Ind. Eng. Chem. Res.(27): 41-45. (1988a). Jayadeokar, S. S. and M. M. Sharma “Ion Exchange Resin Catalysed Etherification of Ethylene and Propylene Glycols with Isobutylene.” Reactive Polymers 20: 57- 67. (1993). Jimenez, L. and J. Costa-Lopez “The production of butyl acetate and methanol via reactive and extractive distillation: II process modeling, dynamic simulation and control strategy.” Ind. Eng. Chem. Res. 41: 6735-6744. (2002a). Marcias-Salinas, R. and J. R. Fair “Axial Mixing Effects in Packed Gas-Liquid Contactors.” Ind. Eng. Chem. Res. 41: 3429-3435. (2002b). Montgomery, D. C. "Design and Analysis of experiments". Brisbane, John Wiley & Sons. (1997). Morita, S. and J. M. Smith “Mass Transfer and Contacting Efficiency in a Trickle Bed Reactor.” Ind. Eng. Chem. Fundam. 17(2): 113-119. (1978). Parra, D., J. Tejero, F. Cunill, M. Iborra and J. F. Izquierdo “Kinetics study of MTBE liquid phase syhthesis using C4 olefinic cut.” Chemical Engineeering Science 49(24A): 4563-4578. (1995). Velo, E., L. Puigjaner and F. Recasens “Inhibition by Product in the Liquid Phase Hydration of isobutene to tert-Butyl Alcohol: Kinetics and Equilibrium Studies.” Ind. Eng. Chem. Res. 27: 2224-2231. (1988b). Velo, E., L. Puigjaner and F. Recasens “Intraparticle Mass Transfer in the Liquid Phase Hydration of Isobutene: Effects of Liquid Viscosity and Excess product.” Ind. Eng. Chem. Res. 29: 1485-1492. (1990). Zhang, C. M., A. A. Adesina and M. S. Wainwright “Isobutene hydration in a countercurrent flow fixed bed reactor.” Chem. Eng. and Processing 43: 533-539. (2003). Zheng, Y. and X. Xu “Study on Catalytic Distillation Processes: Part I: Mass Transfer Characteristics in Catalyst Bed within Column.” Trans. I Chem. E. 70: 459. (1992).

Chapter 7: Synthesis of TBA... 361 8. Basket Impeller Column

In this chapter a new catalytic distillation hardware design is proposed, which is expected to offer many advantages over static forms of hardware such as bale packing. The motivation for the new design, a basket impeller column (BIC), came from various aspects of the previous chapters and from the desire to obtain a better process development tool for the experimental investigation of difficult applications of catalytic distillation, such as the hydration of isobutylene. Those applications limited by large external mass transfer limitations and or suffering from reaction/separation rate mismatch.

The BIC combined the advantages of a dual flow tray with those of a rotating basket slurry reactor. The dual flow tray column has yet to be adapted for the purpose of catalytic distillation. It offers the simplest distillation arrangement possible with few internals, no down comers and the largest free open cross sectional area possible. The liquid and vapour pass through the same openings of the tray, upon which a clear liquid and froth build up. The spinning basket slurry reactor was employed in Chapter 4, under three-phase reaction conditions, to evaluate the kinetics of isobutylene hydration in the presence of ethylene glycol. Its application in a column is simply the special case in which the gas distributor is open enough to allow for liquid weeping of liquid to a another reactor/stage below.

The rotating basket reactor was shown to have very good ability to eliminate external mass transfer limitation. The benefits of the rotating basket reactor in the achievement of uniform fluid composition and intrinsic kinetics studies are well espoused by Carberry (1976). The stirring action prevents bubble coalescence and the fixed bed of catalyst within the basket, experiences high slip velocities. This ability would allow for the high liquid side resistances observed in the CFBR study of Chapter 7 to be reduced without the need of operating at a particular liquid flow rate. If holdup, mixing and liquid flow can be made less dependent on each other then there is a better chance of matching separation and reaction rates.

Chapter 8: Basket Impeller Column 362 The current forms of catalytic packing having relatively poor mass transfer capability are heavily compensated with non-reactive stripping and rectifying sections. The ability to achieve a higher level of separation within the reactive zone could be utilised in improving selectivity and combating inhibition effects. Decoupling of mass transfer artefacts in catalytic distillation and CFBR countercurrent operation would allow for separation effects to be better characterised. External mass transfer limitation may become even more significant should the current isobutylene hydration system be pressurised and two liquid phases be formed. A basket reactor would allow for the two phases to be dispersed using rapid agitation. Such an approach was successfully employed by Gupta and Douglas (1967) and Ihm et al (1988) for the purpose of determination of the kinetics of isobutylene hydration under pressurised conditions within stirred slurry reactors. Further dualflow columns have already been successfully used for heterogeneous systems including that of ethyl-acetate and water, Furzer (2000).

Static systems such as bale packing are reliant on gravity and vapour flow for wetting and penetration. Hence there is not much possibility for mixing and renewal. In the packed type systems only a maximum of about 20 to 25 vol% of the bed can be packed with catalyst. However the catalyst must be oversupplied to allow for deactivation. Consequently, liquid maldistribution and poor gas-liquid contacting may significantly downgrade the efficiency of the reaction zone. Krishna (2002) has recently highlighted some of these challenges and proposed the possible de-construction of catalytic distillation unit to a distillation column with external side reactors or pump arounds. In

Chapter 5 it was observed that there was a minimum in Per,L which corresponded to maximum penetration of the fibreglass bags for bale packing and increased backmixing. Thus a large proportion of liquid flow occurs along the wall of the packing by film flow. However the moving baskets of a basket reactor force the liquid to pass through the catalyst bed. Thus, a severely limiting reagent such as isobutylene can be replenished and an inhibiting product such as TBA can be more effectively removed.

Bale packing is difficult to characterise in terms of variables, which are easily adjusted to allow for different liquid holdup and different throughput. The only opportunity for improved mass transfer or holdup is to add more cloth and use more catalyst, which in turn decreases bed porosity. A BIC type arrangement should allow for holdup to be

Chapter 8: Basket Impeller Column 363 manipulated through rate of rotation and or percentage open area of the tray. These are more readily quantified and manipulated offering quicker and easier process development.

The strength of these motivations and considerations led to the design and fabrication of a single stage evaluation rig, Safinski, Adesina and Young (2003). Information relating to fluid dynamics obtained allowed for the design and fabrication of a functional multistage column, Safinski and Adesina (2005).

The single stage BIC design was based upon dualflow column technology as available in literature. The aim of the study undertaken is to determine the feasibility and the potential range of operation of the new catalytic distillation reactor design. Further, to characterise liquid holdup, froth height and gas holdup, pressure drop and power required in terms of significant variables such as stirring rate, fluid rates, and hardware parameters such as hole size and tray open are and basket size. The variables are first screened using an experimental design study before a deeper matrix design study is reported and practical correlations developed are proposed.

The multistage BIC design is based upon the findings of the single stage study. The aim of the multistage column study is to determine the effectiveness of the proposed system in terms of separation and reaction. With this aim in mind two case studies are discussed. The first focuses upon separation and explores overall column efficiency with an ethanol/TBA test system. The second employs the dehydration of TBA within a reactive separation setup. The order of volatility and the difference in volatility of this system leads to a simple separation when considering dehydration as opposed to hydration, considered in Chapters 4 to 7.

8.1 Literature Relevant to BIC Development An understanding of the dualflow column was sought in order to make a very preliminary judgement of its possible combination with the rotating basket reactor. Further, this understanding was used to make an informed choice of range and selection of manipulated variables. The topics of weeping and application of stirring were also covered, with aim of adding confidence and depth to the analysis and discussion.

Chapter 8: Basket Impeller Column 364

8.1-1 Dual-Flow Columns Dual-flow trays are well known in distillation operations as efficient vapour-liquid contacting devices since they permit countercurrent flow of liquid and gas through the same tray holes. Radial concentration gradient in both liquid and vapour phase is generally negligible because fluid can flow throughout the entire cross sectional area of the tray. Relative to conventional distillation with downcomers there is a paucity of information in open literature on the design of dualflow trays. In one of the earliest studies Myers (1958) observed that tray spacing, percentage open area and hole size have significant effect upon efficiency, loading and flooding points. It was found that increasing tray spacing improved capacity and efficiency although the latter dropped with increased open area. Loading and flooding limits also increased with hole size and open area. Myers worked with a 1” jacketed glass column and used hole diameters of 1/16” to ¼”, percentage open areas of 19 to 35 % and tray spacings of 2 to 3 column diameters. More recent work has confirmed these general findings and further defined the operating loading/flooding window, Takahashi (1986).

Fractionation Research Inc. has conducted extensive commercial research in dualflow column design and operation. They employed a 1.2 m column, using several hydrocarbon systems including n-heptane/cyclohexane, C4 stream components, plus water/alcohol and water/steam systems. Their findings were first described qualitatively by Yanagi (1990) and have recently been documented by Garcia and Fair (2002). Yanagi (1990) observed that industrially dualflow trays offer several advantages. They can be designed for lower pressure drops than cross flow trays. Versions having large hole diameters are well suited to fouling services. Dualflow trays are relatively inexpensive, in terms of both material cost and installation cost. They can be readily used to de-bottleneck columns that suffer capacity limitations. The disadvantages include having lower efficiencies than crossflow trays. Further having a relatively narrow operating range, which results in their turndown ration being low.

Yanagi (1990) described the hydraulics of the column suggesting three operating boundaries:

Chapter 8: Basket Impeller Column 365 • Lower Operating Limit: the condition where the rising vapour applies just enough resistance to the falling liquid so that the latter will not rain through the perforations but be partially held up on each tray. Below this loading point there is insufficient vapour liquid contacting to gain separation. Thus unlike packed columns tray columns require operation above the loading point.

• Upper Operating Limit: as the vapour rate is raised first froth and then spray forms and the efficiency of the column increases. When the spray height reaches the height of tray spacing there is a drop in efficiency. The locus of these efficiency maxima in vapour and liquid velocity for a tray of particular spacing, open area and hole size defines this upper operating limit. The upper operating limit tends to lie in the range of 75 to 90 % of the maximum throughput defined by flooding.

• Flooding: With further increase in vapour rate the entrainment will exceed the downward flowing liquid, causing liquid to accumulate in the inter-tray space and push the column into hydraulic flood. The symptoms of flooding are similar to those for packed columns, excessive column pressure drop, entrainment into the column overhead, deterioration of column efficiency to the point of little or no separation.

The operating and flooding limits are both a function of tray spacing. Thus the peak efficiency attainable also increases with tray spacing as found by Myers (1958). As weeping is a normal operating characteristic of dual flow columns it cannot be used as an indicator of the vapour rate required to gain satisfactory performance. Instead liquid holdup is commonly used. At a sufficiently high vapour rate a pulsating liquid seal is established. Once this seal is established the tray efficiency increases with each incremental increase in vapour rate till the upper operating limit is reached. The effects of percent open area and hole size on flooding are apparent, but not well understood. Smaller open areas flood more readily with increased vapour velocity as they promote jetting which increases liquid entrainment. Larger holes give stronger longer jets thus they suffer from flooding at marginally lower vapour velocities (6% capacity) than smaller holes, Garcia and Fair (2002).

Chapter 8: Basket Impeller Column 366 Yanagi (1990) observed that any single hole of a dualflow tray does not pass liquid and vapour simultaneously. Each is either passing vapour, liquid or neither thus each is in one of three possible states. The vapour can be passed as either jets or bubbles. Due to the fluctuating nature of the flow at the tray, pressure drop across a given tray can also fluctuate. However, these effects tend to average out over a number of trays within a column, suggesting that each tray tends to act independently of the others. This fluctuation is also key to the self-cleaning ability of dualflow columns.

Garcia and Fair (2002) proposed that for either side of the peak efficiency the mechanism of loss of efficiency is that of entrainment, albeit occurring in two different forms. The upper loss of efficiency was attributed to the standard form of entrainment, where loss of efficiency occurs through recycle of liquid carried in the vapour back to the tray above. The lower end loss of efficiency results due to entrainment in reverse or bypass of liquid of the tray to the next below. They applied the approach of the Colburn (1936) for characterising the affect of entrainment given by:

E 1 = (8.01) E E Ψ p 1+ p ()1− Ψ

where Ψ is the fraction of liquid entrained and Ep is the peak efficiency. Similarly defined parameters were used for entrainment in reverse. For a given spacing, values of

Ep tend to be higher for trays of lower percentage open area and naturally occur at lower vapour velocities. Garcia and Fair (2002) suggested that for industrial applications dualflow columns of 15-20 % open area, 12.7 mm holes and 0.5 to 0.7 m tray spacing are an adequate compromise of efficiency, pressure drop and operating range. The range of superficial velocities for such trays was that of 0.25 to 2.0 m.s-1. Studies concerning dualflow trays tend to place great emphasis on high capacity operation.

Within the normal operating range as the vapour velocity increases the tray hydraulics go through different regimes such as frothing, transition and incomplete spray. Gas- liquid interfacial area decreases from froth to transition regimes Miyahara (1990). Xu, Afcan and Chuang (1994) have obtained useful correlations to predict the number of gas

Chapter 8: Basket Impeller Column 367 and liquid side transfer units and thus mass transfer coefficients for dualflow columns. They studied methanol/water and methanol/2-propanol distillation in a 0.3 m column and provided valuable insight into the mass transfer involved. They also validated the correlation of Mahendru and Hackle (1979) for froth porosity given by:

0.2 2 − ⎡UGGρ ⎤ HGfroth. =−1.0 0.0946 ⎢ ⎥ (8.02) ⎣ ghLLρ ⎦ where:

hhL=− froth(1 H G. froth ) (8.03)

and adapted that for equivalent clear liquid height, hL, (the height of liquid obtained if the vapour is removed from the froth) using their own regression parameters obtaining:

1.75 ⎡ 0.5 ⎤ n ⎛⎞ρG ()LMLG⎢ U ⎜⎟⎥ ⎢ ⎝⎠ρair ⎥ h = 0.006 ⎣ ⎦ (8.04) L 1.9 0.42 ⎛⎞Ah ⎛⎞t ρL ⎜⎟⎜⎟ ⎝⎠Ada ⎝⎠ where:

−0.25 ⎛⎞A n = 0.3162⎜⎟h (8.05) ⎝⎠Aa

and where Ah is the hole area and Aa is the column active or cross sectional area. Equation (8.04) can be used to correlate for pressure drop as suggested by Garcia and Fair (2002).

Furzer (2000) has extended dualflow trays to complex azeotropic mixtures both heterogenous and homogeneous. He has attempted to characterise the dynamics and identified the positive weeping of dualflow trays as having a strong interaction with froth height, Furzer (2001). He noted that the oscillation or fluctuation of hole state produces mixing and that the oscillation increases with decreased column diameter.

Chapter 8: Basket Impeller Column 368 Furzer (2001) used a pilot scale column of 6 m height, with 0.15 m diameter, 0.32 m tray spacing, 8 mm hole diameter and 20 % open area.

An understanding of weeping at the hole or orifice may improve the design of dualflow columns. The next sub-section looks at studies conducted specifically looking at the phenomena associated with weeping, both for its enhancement and prevention.

8.1-2 Weeping and Bubble Formation at a Submerged Orifice Akagi et al (1987) used high-speed photography to demonstrate that a cycle of phenomena occurred at the orifice including bubble formation, detachment and liquid weeping. They tested weeping at a single submerged orifice of a tray dividing a gas chamber of variable volume and a certain maintained level of liquid. They measured flow rates and pressure and found that the gas chamber pressure rose with accumulation of gas and then the bubble formed at the orifice. The enlargement and detachment of the bubble occur in a moment and buoyancy helped form and lift the bubble. The chamber pressure decreases excessively and the liquid weeps immediately after the detachment of the bubble. For a single orifice they observed two regions in weeping rate behaviour. In the first region weeping increases steadily with gas flowrate and reached a maximum. In the following second region, weeping rate decreases until total cessation.

Peng, Yang and Fan (2002) made a number of observations relating to weeping and the investigation of weeping. They suggested that:

• Deviation of the results of various studies conducted may be due to tray material wettability which can influence bubble formation and showed that more wettable stainless steel trays behave differently to plastic ones. • For water-based test systems, situations other than that of single bubble formation can occur, such as cluster formation, at almost any vapour flowrate due to water’s abnormally high surface tension and the strong influence of water contaminants upon bubbling. Measures of weeping rates in water can have as much as 40 % standard deviation. Water’s high surface tension allows for fast reestablishment of the gas liquid interface. For 1/8” orifice single bubbling could not be maintained beyond 8 cm.s-1.

Chapter 8: Basket Impeller Column 369 • During the bridging period of pressure buildup before bubble formation some weeping occurs by a different mechanism to the bulk of weeping. • The single orifice situation neglects the reality of strong mixing or circulation that may result from a group of orifices interacting with one another. • Columns 4 to 12 times larger than the average bubble diameter may still be subject to wall effects, and that larger column diameters experience higher rates of weeping and less stable bubbles. • Operating in the jetting regime may not be sufficient to prevent weeping due to circulation effects. That cross flow above the orifice plate leads to moments of jet instability and increased pressure directly above the orifice that induce weeping. As larger orifices provide a higher gas momentum force to maintain jet stability, they are subject to reduced weeping rates in the jetting regime. The jetting regime for a 1/16, 1/8, and 1/4 “ orifice was found to occur at orifice velocities beyond 10 m.s-1. • Further that larger orifices lead to higher weeping rates. An increase from 1/8 to 1/4” for an aqueous system results in a weeping rate increase of 5 times.

Zhang and Tan (2003) attempted to decouple the effect of liquid circulation upon multiple orifice weeping through the use of baffled orifices. It is difficult to judge whether they achieved the desired objective of eliminating recirculation or created a new system all together. Their results suggested that weeping rates were higher in the absence of circulation, which is contrary to the findings of Peng, Yang and Fan (2002). It is possible that they created a hybrid of the form of orifice flow modification as proposed by Billet (2001), called Dualflex trays. In the case of Dualflex trays mobile valve plates and top weirs positioned above the perforations of the tray create a guided dualflow action which promotes mass transfer and eliminates the degree of oscillation observed with straight orifice use. Zhang and Tan (2003) additionally demonstrated that weeping increased with the head of liquid, but became stable after a certain critical value.

8.1-3 Rotation and Distillation Rotation has been previously used with distillation and two example of its application are high gravity distillation and the slurry distillation using a spinning cone column.

Chapter 8: Basket Impeller Column 370 In high gravity distillation the required separation is performed radially in a high speed (750 to 1000 rpm), high centrifugal force rotating packed bed, Kelleher and Fair (1996). The bed is a single cylinder and the system relies on perfect balancing and advanced gas seals. These columns have been commercialised for use on oil platforms, where space is a major consideration.

The spinning cone column is used primarily in the food industry to strip slurries of volatile components such as flavours or alcohol when both resulting slurry and distillate are of value, Wright and Pyle (1996). It uses a series of stationary cones, mounted on the wall side and another series of mobile cones, mounted on a central rotating shaft. Together these allow for a finely spread slurry to pass through the column by combined centrifugal and gravity action, achieving very high mass transfer coefficients. These examples although different in operational principle to the proposed design, demonstrate that enhanced separation and or particular hardware requirements can offset the additional power input of rotation.

Chapter 8: Basket Impeller Column 371 8.2 BIC Single Stage Evaluation The approach to evaluation of feasibility of the proposed new system is that of constructing a single stage of the proposed column design, which could be characterised in terms of fluid capacity. Of primary interest is the effect of design parameter basket size, sieve plate open area, and hole size upon gas and liquid holdup. Rotational speed and gas and liquid rates were also varied as operation parameters. The simple experimental rig used is depicted in Figure 8.01 (a).

8.2-1 Single Stage Apparatus The single stage comprised of test plate, tray space column equipped with baffles, liquid distributor and calibrated collector, shaft and basket. The tray space column and liquid collector were connected by a flanged union, which held the test plate between rubber gaskets as depicted in Figure 8.01 (a). The tray space column itself was fabricated from clear plexiglass to permit the determination of froth height and visual monitoring of other hydrodynamic phenomena. The tray space column was equipped with four equally spaced square baffles of 5 mm width and 70 mm height located immediately above the test tray. The tray space had a height of 0.5 m. A large diameter of 0.15 m was chosen to ensure that:

• The design was not only feasible at a very small scale • Wall effects were minimised, given the diameter of the holes tested was 4 and 6 mm this gave an aspect ratio of 31:1. • The equilateral triangular pattern of the tray perforation could be maintained and an even well characterised distribution achieved. As the column diameter decreases this pattern becomes difficult to maintain. • To obtain measurable difference in liquid holdup.

Details of the weir type liquid distributor are given in Figure 8.01 (b). The simple design and triangular cuts of the individual weirs ensured good distribution at relatively low flowrates. The design was based upon that of the Norton Co. weir type distributor, Sinnott (1997).

Chapter 8: Basket Impeller Column 372 The column was equipped with a scaled liquid level sight, which allowed for calibration and subtraction of the liquid in the collector from the overall measurement.

The impeller baskets, Figure 8.01 (c), were fabricated by soft soldering 304 SS Melwire mesh of 315 μm aperture and 180 μm wire diameter. The baskets were of 120 mm diameter, 10 mm arm width and three different heights: 10, 30 and 50 mm. They were maintained at 5 mm above the test tray. Glass spheres of 850 to 1180 μm were used as inert particles representing a catalyst such as Amberlyst 15. Each of the six, impeller arms was balanced before they were glued in place around a central plastic structure with epoxy resin. At this point it was recognised that a multi-arm basket would give a short liquid path length through the catalyst bed contained within. That a greater number of longer arms would give a higher inertia of rotation and would be more suited to stable rpm operation than dynamic varying rpm operation.

The sieve trays were made of t = 2 mm thick PVC plate and were drilled in an equilateral triangular pattern. The drilling resulted in very clean-cut straight holes. Test plates of dh = 4 and 6 mm and percentage open area (%OA) of 10 and 20 % were fabricated. Values were chosen at the lower end of the standard dualflow column design, given the expectation that forced circulation would enhance weeping. The hole- pitch, lp, for the equilateral triangular pattern was estimated using, Sinnott (1997):

2 ⎛⎞Ad⎛⎞ hh= 0.9 (8.06) ⎜⎟ ⎜⎟ ⎝⎠Alap⎝⎠

Hole pitch values and the corresponding percentage open areas are given in Table 8.01.

Table 8.01 Estimate of hole pitch, lp, required to achieve certain % Open Area. %OA: 10 20 dp (mm) Pitch (mm) 4 12 8 6 18 13

The sieve trays were held between two flanges with 12.5mm bolts and the column was sealed with two 2mm rubber gaskets. A digital stirrer Heidolph RZR 2021 was applied

Chapter 8: Basket Impeller Column 373 to a 5 mm shaft, which was held in place by two teflon guides, which allowed for travel in the vertical direction. A gravimetric method was used to determine liquid holdup via a continuous, fast response, 1 g resolution, Wedderburn, Precisa 40,000G IPG5 scale. The column, stirrer and balance were mounted within a square tube aluminium frame, which was levelled and weighed down with two 20 kg masses for stability. The column could be supported or lowered such that its entire mass rested upon the balance.

8.2-2 Single Stage Procedure The holdup in the liquid phase was measured gravimetrically. Air and water at atmospheric pressure and ambient temperature were used as fluids. Water flow rate was measured with a calibrated Dwyer 20 gph flow meter. Air flow rate was regulated with a Whitey valve which was pre-calibrated. Flow rates were chosen to simulate distillation conditions with air superficial velocities between 0.2 and 0.9 m/s. The gravimetric method of capacitance determination involved establishing a stable gas and liquid flow rate at a given impeller speed followed by stabilisation of the bottom liquid level with the drain valve. The bottom liquid level, mass, froth height, and pressure drop were then recorded. Measurements were taken at least three times over the run period (post-stabilisation) in order to obtain an average value.

Chapter 8: Basket Impeller Column 374

[01] RPM

[02]

[03] [07] [04] [05] [06] P

[09] [08]

MASS [10]

D 12 mm

40 mm

B1 (b) (c) B2

B3

Figure 8.01 Basket Impeller Column: (a) Holdup Determination and Feasibility Rig, (b) liquid distributor design & (c) basket impeller design

Key to Figure 8.01 Basket Impeller Column: (a) Holdup Determination and Feasibility Rig

Chapter 8: Basket Impeller Column 375 [01] Digital Stirrer, Heidolph, RZR-2021, 0 to 2000 rpm. [02] Liquid Distributor. [03] Clear Column, Acrylic, 1500 mm od, 3 mm thickness. [04] Basket Impeller, three sizes mounted on 5 mm ss shaft. [05] Baffles. [06] Test Plate. [07] Water Supply, Rotameter, Dwyer, Ratemaster RMC. [08] Air Supply, calibrated valve, Whitey ss ½”. [09] Calibrated Liquid Level Sight, ½” scaled glass tube. [10] Digital Balance, Wedderburn, Precisa 40,000G IPG5.

Basket Impeller Column Single Stage Experimental Procedure The measured values were overall unit mass, froth height and pressure drop. The following method was employed: 1. Zero digital balance [10]. 2. Assemble the BIC unit with the desired sieve plate [05] and basket [06]. Place within experiment frame, centralise and attach stirrer [01]. 3. Take down the unit mass. (Empty) 4. Set the desired RPM. 5. Take down the unit mass. (Vibration/balance) 6. Set desired liquid flow rate with valve and rotameter [07]. 7. Adjust liquid level in the liquid collector/calibrated volume [09]. 8. Take down unit mass and liquid level. (Wet basket/distributor) 9. Set desired air flow rate using pre-calibrated valve[08]. Make sure that the applied utility air pressure is consistent with previous runs. (600 kPa) 10. Allow sufficient time for the column to stabilise. Take froth height measurements and pressure drop measurements regularly to help determine steady state. 11. Take down the unit mass and liquid level. (Aerated Value) 12. Measure the froth height and pressure drop. Shut down the unit drain and establish new set point.

Chapter 8: Basket Impeller Column 376 8.2-3 Experimental Design Two single stage studies were conducted, Safinski, Adesina and Young (2003). The aim of the first study is to determine the feasibility of the proposed BIC design and distinguish those variables and interactions, which could be considered significant. The aim of the second study is to develop a correlation, which may be used to describe the main measures of liquid and gas holdup, froth height and pressure drop. A third supplementary study was conducted with the aim of characterising the power associated with stirring the gasified froth.

The first single stage plate design study was conducted according to a full factorial experimental design of five factors, which included gas and liquid volumetric flowrates

(FG and FL), Ω (rpm), plate hole diameter (dh) and plate percentage open area (%OA). The variables were varied as shown in Table 8.02. A graphical method of normality plot construction was employed to determine significant factors and interactions, instead of the use of ANOVA analysis as employed in Chapter 7.

Table 8.02 Single Stage Tray study full factorial experimental design Factor Definition Low (-1) High (+1) Units -1 A QG 242 495 L.min -1 B QL 0.126 0.63 L.min C Ω (rpm) 100 300 rpm D dh 4 6 mm E %OA 10 20 -

Additional runs exploring each variable in turn for the purpose of trial and observation also supplemented these design runs. For the purpose of the second study the hole diameter dh was fixed at 4mm and the tray %OA at 10 % based upon the results of the first study. Fluid flowrates, Ω (rpm) and basket size previously not studied were varied according to Table 8.03.

Table 8.03 Single Stage Extended Capacity Study Experimental Variables Variable Values h basket (mm) 10, 30, 50 ReG 265, 613, 961, 1309, 1657, 2005, 2353 ReL 5, 12, 25 Ω (rpm) 50, 75, 100, 125, 150, 175, 200, 225, 250, 275, 300

Chapter 8: Basket Impeller Column 377

The third short study focused on measurement of power consumption, which was made possible through the acquisition of a digital stirrer capable of measuring torque, Heidolph RZR 2102. Preliminary attempts had been made with voltage and current measurement but the results were considered only indicative, as motor efficiency was difficult to estimate. The Heidolph RZR 2102 could be zeroed and only the torque needed to rotate the stirrer under different column conditions was measured at a resolution of ± 0.1 N.m. The measurements of torque were used in conjunction with the knowledge of RPM to calculate power, P (W), by the formula:

PT= ω (8.07) where ω is angular velocity and T is torque. Torque measurements were made at ambient temperature and pressure. Air flowrate was varied to give the range ReG 316 to

2280. Water flowrates were varied giving ReL 11 to 36. The speed of rotation was varied in the range RPM 0 to 500 giving the range ReI 250 to 2000.

Chapter 8: Basket Impeller Column 378 8.2-4 Single Stage Sieve Tray Study, Results and Discussion The initial study indicated that a feasible design for the BIC column is possible and attainable. Preliminary runs were conducted to establish the loading and flooding points for the sieve plates. Fluid flow rates were therefore chosen to be above the loading point. The stirrer was seen to have a pronounced effect upon bubble size, froth behaviour and the extent of the clear liquid phase.

Baffles, which were employed in the design to encourage recirculation, were observed to have a marked impact especially as Ω was increased to 500 rpm. The relatively small baffles diverted the wall liquid and froth back towards the centre.

Following the experimental design of Table 8.02, the capacity (g) was measured and the results are reported in Table 8.04.

Table 8.04 BIC Initial Plate Design Study Results Run A B C D E Capacity QG QL Ω (rpm) dh %OA (g) 1 -1 -1 -1 -1 -1 193 2 1 -1 -1 -1 -1 497 3 -1 1 -1 -1 -1 455 4 1 1 -1 -1 -1 1039 5 -1 -1 1 -1 -1 39 6 1 -1 1 -1 -1 154 7 -1 1 1 -1 -1 73 8 1 1 1 -1 -1 327 9 -1 -1 -1 1 -1 251 10 1 -1 -1 1 -1 509 11 -1 1 -1 1 -1 231 12 1 1 -1 1 -1 776 13 -1 -1 1 1 -1 116 14 1 -1 1 1 -1 261 15 -1 1 1 1 -1 115 16 1 1 1 1 -1 254 17 -1 -1 -1 -1 1 94 18 1 -1 -1 -1 1 276 19 -1 1 -1 -1 1 49 20 1 1 -1 -1 1 397 21 -1 -1 1 -1 1 54 22 1 -1 1 -1 1 71 23 -1 1 1 -1 1 83 24 1 1 1 -1 1 135 25 -1 -1 -1 1 1 0 26 1 -1 -1 1 1 195 27 -1 1 -1 1 1 0 28 1 1 -1 1 1 64

Chapter 8: Basket Impeller Column 379 29 -1 -1 1 1 1 0 30 1 -1 1 1 1 0 31 -1 1 1 1 1 4 32 1 1 1 1 1 24

The method of analysis of the experimental design was different to that used in Chapter 7 for CFBR and CED experimental investigations. Instead of resorting to a fractional factorial design a simplified method of analysis was adopted to minimise the number of runs. The method detailed in Montgomery is particularly suitable when applied to a design of four or more factors. It involves calculating the effects in the usual manner. Sorting these in ascending order while maintaining information pertaining to their identity and calculating their normal probability defined:

(k − 0.5) P = (8.08) k n where k is the order and n is the total number of effects. A straight line is next constructed through typically the points of probabilities 0.25 and 0.75. Factors for which the effects plotted in this manner do not lie upon the line of normality can be considered to be significant. The test works well with a large number of factors as the fit and range of the line are more easily determined and defined.

The results of the analysis of the experimental design were captured in the normality plot of Figure 8.02, which indicated that the most significant variables affecting liquid holdup were those of gas flow rate, %OA and the impeller speed. Liquid flow rate had only slight significance while variation in hole size, was inconsequential to the results. There was, however, a significant positive interaction between the impeller speed and %OA. The plate with 10 %OA offered greater liquid retention than the 20%OA design, however, as the impeller speed increased the difference between the two plates diminished. Although hole size appeared not to influence liquid holdup, with smaller hole size the liquid retained displayed better frothing probably due to higher interfacial area of smaller bubbles formed and hence good backmixing.

Chapter 8: Basket Impeller Column 380 120

100 ReG RPM & %OA 80

60 % j P 40

20

RPM 0 %OA

-300 -200 -100 0 100 200 300 Effect

Figure 8.02 Normality plot for the experimental design of the first study, indicating significant factors and interactions

A wide variability in tray liquid capacity (g) could be achieved with a plate with smaller %OA over the same range of gas superficial mass velocity. This feature will be advantageous in reactive distillation since a wider range of Damköhler number will be admissible. On the basis of these results, the optimum plate with 4 mm hole diameter and 10 %OA was chosen for further investigation

Chapter 8: Basket Impeller Column 381 8.2-5 Single Stage Fluid Dynamics Study An extended study on capacity was carried out for the plate with 4 mm hole diameter and 10 %OA. This study focused on the effect of impeller size upon liquid holdup, gas holdup within the froth, pressure drop and power requirements. Each of the measures: liquid phase holdup, gas phase froth holdup, froth height and pressure drop is discussed in turn and simple yet effective correlations are proposed.

8.2-5-1 Liquid Phase Holdup

The sensitivity of the tray liquid holdup towards impeller size was explored in the second study in order to project the potential maximum catalyst loading, which the baskets may offer. Increasing the size of the impeller basket strengthened its mixing and shearing action. In general liquid holdup diminished almost exponentially with increase in impeller speed, as evident in the case of the largest basket (Basket 3), Figure 8.03.

0.20 0.18 Basket 1. Basket 2. 0.16 Basket 3. 0.14 0.12

L 0.10 H 0.08 0.06 0.04 0.02 0.00 0 50 100 150 200 250 300 RPM

Figure 8.03 Liquid Holdup with Basket size and RPM at ReG 960, and ReL 12

Chapter 8: Basket Impeller Column 382 The liquid holdup refers to the tray space volume defined by the tray and liquid distributor of an area of 0.018 m2 and height 0.4 m, giving 0.0071 m3 or 7 L. Thus for a 10 %OA tray the froth is reduced to approximately 100 g of water at Ω = 250 rpm. The basket impeller column is a special case of the dualflow column with heightened weeping due to forced circulation. As rate of stirring was increased the frequency of jet and bubble disruption increased giving higher rates of weeping and lowering liquid holdup or retention.

The effect of Ω upon froth behaviour was captured in a series of photographs taken of the BIC during experimentation Figure 8.04. With increased Ω the froth can be seen to initial expand, however as the rate of weeping increases the froth is starved of liquid and begins to collapse. It is eventually reduced to a continuous spray between the arms of the basket and the baffles.

Figure 8.04 The effect of RPM. Left to Right RPM 0, 50, 75, 100, 125 and 150 at

ReG 960 and ReL 12

As expected for a countercurrent system the liquid holdup, HL, increased monotonically with ReG. Since both fluids share the same openings in a dual flow tray, as gas flow increases more liquid is retained on the tray. This effect becomes amplified at higher liquid velocities, Figure 8.05.

Chapter 8: Basket Impeller Column 383

0.4

ReL 5

ReL 12 0.3 ReL 25

0.2 L H 0.1

0.0

0 500 1000 1500 2000 2500

ReG

Figure 8.05 HL with ReG at RPM 100 and ReL 12

Again a sample of the observed effect of ReG upon froth was captured in a series of photographs taken during experimentation, Figure 8.06 which corresponds to the plot of Figure 8.05.

Figure 8.06 The effect of ReG. Left ot right ReG 265, 613, 960, 1309 and 1656 at

RPM 100 and ReL 12

Chapter 8: Basket Impeller Column 384 As only a single stage was used a full flooding and loading envelope have not been defined. However the full possible stage froth height was achieved for each liquid flow rate by incremental increase in the gas flow rate to the onset of flooding. This was typically achieved in all cases with 2000 < ReG < 2500.

The liquid holdup, HL, at various combinations of gas and liquid flow rates was correlated with the expression:

0.8609 ⎛ h ⎞ −4 1.2693 0.5254 −0.6749 ⎜ I ⎟ L ×= G L ReReRe101.183H I ⎜ ⎟ (8.09) ⎝ w I ⎠

Figure 8.07 displayed the parity plot between predicted and observed values of HL at 95% confidence level from more than 250 runs. The error of prediction of the model of

HL was ± 0.005. The model displayed some tendency to overestimate the lower range values of holdup.

0.5

0.4

0.3

Correlated 0.2 L H

0.1

0.0 0.0 0.1 0.2 0.3 0.4 0.5

HL Observed

Figure 8.07 Parity plot for Equation (8.09)

Chapter 8: Basket Impeller Column 385 8.2-5-2 Gas Phase Holdup

The total gas holdup, HG, is made up of the fraction in the froth and the disengaging height immediately above the frothing liquid, thus:

HHGGFrothGDH= ,,+ H (8.10) Additionally:

HHGL+ =1 (8.11)

It is the gas holdup in the froth that is the most critically affected by changes in the impeller speed, hole size and percent open area. It is the gas held by the froth that undergoes the greatest extent of mass transfer. The total froth volume was estimated from the measured froth height. As a result:

VolG,Froth HG = (8.12) VolStage where

VolG, Froth= Vol Froth−− Vol L Vol Basket (8.13)

The froth gas holdup, HG, Froth, was essentially determined by measuring froth height. As shown in Figure 8.08 (a), HG, Froth dropped in the same manner as the liquid holdup attaining a flat profile at high RPM. The decrease in HG, Froth was probably due to the gas displacement into the disengaging portion of the column due to reduced froth volume occasioned by the enhanced shearing and weeping of liquid through the holes. Not surprisingly HG,Froth increased with gas flowrate (ReG) almost linearly. Interestingly the slopes of the HL vs. ReG and HG, Froth vs ReG plots were of about the same order of

Chapter 8: Basket Impeller Column 386 0.35 Basket 1. 0.30 Basket 2. Basket 3. 0.25

0.20 G,Froth H 0.15

0.10

0.05 0 100 200 300 RPM

0.5

0.4

0.3

0.2 G,Froth H Re 5 0.1 L

ReL 12 0.0 ReL 25

400 800 1200 1600 2000 2400 2800

ReG

Figure 8.08 HG with (a) RPM and basket size and (b) ReG and ReL

Chapter 8: Basket Impeller Column 387 magnitude (10-4) and this is a reflection of the resistance (geometry and material type) of the plate to competitive passage of both gas and liquid through the same holes.

8

7 Basket 1. Basket 2. 6 Basket 3.

L 5 /H 4

G,Froth 3 H

2

1

0 0 50 100 150 200 250 300 RPM

Figure 8.09 The ratio HGFroth/HL with RPM showing the greater extent of aeration and recirculation of bubbles

Indeed the ratio of HG, Froth to HL, which is a measure of aeration increased with impeller speed as may be seen in Figure 8.09. Higher rotational speeds result in faster rates of weeping of liquid through the tray holes, but promoted bubble back mixing. The bubbles swarmed around, tracing out a spiral in the direction of impeller rotation. Increased flow rate produced larger bubbles with stronger upward flow rates, which were able to break free of the liquid in a jetting manner. Increased shaft rotation produced a thicker, clear liquid zone and although backmixing was increased its depth and penetration into the froth was reduced. The froth surface was obviously highly agitated but there was no evidence of coning or central cavitation, hence it is unlikely that froth was spread up against the column wall at very high impeller speeds. The plot

Chapter 8: Basket Impeller Column 388 also revealed that at low speeds (rpm<100) the basket size did not have a significant effect on the degree of aeration. However at higher speeds the aeration was improved with increased basket size.

This interplay of hydrodynamic factors was correlated with gas hold up as given below:

−7.9983 A −0.3685⎛ h I ⎞ HG = 0.0177ReGReL ⎜ ⎟ + 0.5308 ⎝ w I ⎠ (8.14) −5 []×−= III − 0.355)/w(hRe101.7A

The resulting parity plot given in Figure 8.10 lends credence to the reliability of this correlation. The parity plot showed that the model given by Equation (8.14) tended to over

0.7

0.6

0.5

0.4 Correlated 0.3

G,Froth 0.2 H 0.1

0.0 0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7

HG,Froth Observed

Figure 8.10 Parity plot for Equation (8.14)

estimate HG,Froth at small values of holdup. However estimate uncertainty for HG according to the model was quite acceptable at ± 0.0018 vol/vol.

Chapter 8: Basket Impeller Column 389 8.2-5-3 Froth Height

Froth height, hf (mm), is an important parameter as conventionally the column is considered to be flooded if the froth height reaches the tray above. Thus during the design phase the tray spacing needs to be sized appropriately using knowledge of froth height. It was correlated with the hydraulic variables to yield:

6.5215 ⎛ h ⎞ 0.9715 0.4028 −0.5039 ⎜ I ⎟ hf = 0.7266ReG L ReRe I ⎜ ⎟ (8.15) ⎝ w I ⎠

Figure 8.11 showed that Equation (8.15) is an adequate representation of the large volume of data collected with froth height accurate to within ± 7 mm.

500

400

300 Correlated 200 Froth h 100

0 0 100 200 300 400 500

hFroth Observed

Figure 8.11 Parity plot for Equation (8.15)

Chapter 8: Basket Impeller Column 390 8.2-5-4 Pressure Drop

Pressure drop seemed to decrease with increased impeller speed similar to the trend seen in the liquid hold-up dependency on the same flow variables, Figure 8.12 (a). As expected, however, increased gas flow rate, ReG, increased system pressure drop as shown in Figure 8.12 (b).

1500 1400 Basket 3 ReG 961, ReL12 1300 1200 1100 P (Pa) Δ 1000 900 800 700 0 50 100 150 200 250 300 RPM

3500

3000

2500

2000

P (Pa) 1500 Δ

1000 ReL 5 Re 12 500 L ReL 25 0 500 1000 1500 2000 2500

ReG

Figure 8.12 Pressure drop with (a) RPM and (b) ReG

Chapter 8: Basket Impeller Column 391 The dependency on ReL is, however, not as strong as that for ReG. This qualitative observation is reflected in the exponent with respect to ReI, ReL, ReG and the impeller aspect ratio (hI/wI) as given by: 0.05157 ⎛ h ⎞ 1.1988 0.1121 −0.3236 ⎜ I ⎟ ΔP = 0.1635ReG L ReRe I ⎜ ⎟ (8.16) ⎝ w I ⎠ with an associated uncertainty in ΔP of ± 81.04 Pa. The parity plot of Figure 8.13 showed a no particular trends and was generally random giving support to the adequacy of Equation (8.16) for ΔP.

350

300

250

200

150 P Correlated Δ 100

50

0 0 50 100 150 200 250 300 350

ΔP Observed

Figure 8.13 Parity plot for Equation (8.16)

Pressure drop decreased with increased RPM and it may be suggested that some of the energy given to stirring could be recovered in a reduction in the energy required for boilup. The complete balance will be very chemical system dependent, but will be something to consider when the complete column is operated.

Chapter 8: Basket Impeller Column 392 8.2-5-5 Power Requirement

Power consumption versus process gains or alternative process design will become a major consideration in BIC application. The addition of stirring to separation columns has been proven viable in the field of extraction, slurry distillation and high gravity distillation. In extraction stirring is used to assist redispersion after each stage section. Columns such as the Scheibel and rotating disc column have been developed and commercialised. For the purpose of slurry distillations such as alcohol removal the spinning basket column has been developed, Wright (1996). For distillation under highly confined space as oil rigs and ships high gravity distillation has been designed, Kelleher (1996). Stirring need not be power intensive if the unit is continuous, stirring can be maintained at one speed, the speed is not very high, drag is minimised and devices such as fly wheels are used.

As depicted in Figure 8.14 power of stirring increased almost linearly with stirring rate Ω (rpm) for a given combination of fluid flow rates, as may be expected.

1200

1000

800

600 P (W) 400

200 Basket 3 ReL 23 ReG 1354

0 0 100 200 300 400 500 RPM

Figure 8.14 Power P (W) with RPM at ReG 1354 and ReL 23

Chapter 8: Basket Impeller Column 393

The power requirements observed were adequately correlated with the following power law expression:

0.1446 0.0544− 0.0215 1.1357 ⎛⎞hI PGGLI= 0.1160Re Re Re ⎜⎟ (8.17) ⎝⎠wI

A good fit was obtained of R2 = 0.9974. It could be seen from Equation 8.17 that the most influential variable was RPM and that the influence of the fluid Reynolds numbers was generally weak in comparison, but nevertheless interesting as they were considered to be counter-intuitive. There was a very minor drop in power with ReL and a minor rise with ReG, when the opposite was expected. These effects may be due to turbulence. The gas flowrates used were quite high approaching 1 m3/min, and as gas flowrate was increased the occurrence of jets of gas and bubbles increased. Crossing over strong jets may act as a brake, requiring additional energy to maintain speed of rotation. An increase in both ReG and ReL increases liquid phase holdup, but ReG is a stronger influence Equation 7.01. However ReG tended to create froth while ReL created clear liquid. Clear liquid may have helped shield the stirrer from the turbulence of the gas jets.

The power requirements of stirring flat liquid were correlated against the mass of liquid,

M (kg), and ReI. The resultant correlation was found to be:

0.0673 1.2502 PM0 = 0.0771LI Re (8.18)

where ML is mass in kg. The fit obtained was 0.9959. Equation (8.18) was used with Equation (8.09) for holdup under gasified conditions and it was found possible to compare the power requirements of stirring flat and gasified liquid on a mass basis. This type of comparison had been made by Bruijn, vant Riet and Smith (1974). in their extensively referenced work on Rushton turbines They found that the ratio of gassed power, PG, to ungassed power, P0, as a function of operating conditions could be captured by

Chapter 8: Basket Impeller Column 394 −0.25 24−0.20 PFGG⎛⎞⎛⎞RPS dI = 0.10⎜⎟⎜⎟2/3 (8.19) P0 ⎝⎠ RPS. V⎝⎠ gwV

3 3 where V is the volume of liquid (m ), FG the volumetric gas flowrate (m /s), dI the impeller diameter (m), w its blade width and g the acceleration due to gravity.

The standard Rushton turbine system is quite different in geometry to the BIC. Further the flow regimes employed are substantially different, however the same dimensionless numbers as developed for Equation (8.19) can still apply as the liquid flow rate is much smaller than the gas flow rate. The second dimensionless number was modified slightly to include basket height h and basket width was used as w. A non-linear regression was carried out for the data generated for the BIC and the following correlation was obtained:

0.1364 23 −0.0029 PGBIC, ⎛⎞FG ⎛⎞RPS dI h = 0.9538⎜⎟⎜⎟2/3 (8.20) P0,BIC ⎝⎠ RPS. V⎝⎠ gwV

It could be seen from the exponents that the ratio became smaller with increased gas flowrate in this high gas flow turbulent regime. Over the range of the study the ratio remained greater than unity. However it can be seen that with sufficiently high RPM and gas flowrate the value can drop below unity and there will be no energy gained from gasification or increased vapour rates through additional boilup.

The most important variable in terms of power quite obviously remains RPM. The BIC was designed with operation at moderate RPM of 50 to 500 in mind. For a given process an important optimisation in terms of RPM and power will arise. Using the correlation developed will allow for this optimisation.

Chapter 8: Basket Impeller Column 395 8.3 BIC Multistage Case Studies A multistage BIC was designed and fabricated. A cassette type design was adopted with the internals inserted as one complete structure into the column shell. The catalytic distillation column was applied to the separation of an ethanol/TBA mixture and to the dehydration of TBA.

The ethanol/TBA system was chosen for the first separation study. It is more typical of those found in current applications of catalytic distillation, Quitain et al (1999) than conventional test systems, Fair (1987). It exhibits ideal very flat VLE, well suited to efficiency determination. The similarity in physical properties of TBA and ethanol, Sinnott (1997), results in consistent fluid dynamics over the length of the test column.

The dehydration of TBA over Amberlyst 15 to isobutylene and water:

H+ CH COH ⎯⎯→ CH C=CH + H O (8.21) ()332232←⎯⎯ () was chosen for the first reactive application. This system is equilibrium limited in the same manner as the hydration of isobutylene. It is relatively simple and the order of volatility and the difference in volatilities between the products ensures good separation and ease of operation. For these reasons this system has been previously used to assess the effectiveness of new catalytic distillation reactor designs, Gotz et al (2001) and Tuchlenski et al (2001).

8.3-1 Multistage Column Apparatus The new basket impeller dualflow catalytic distillation column setup is shown in Figure 8.15 and detail of the cassette type internals is shown in Figure 8.16. Based upon the experience of the single stage study, dh = 4 mm, %OA of approximately 5 and 10 %, and tray spacing equal to three diameters were employed. The multistage column fabricated is essentially a 9-stage, 1.8 m long polycarbonate column (ID=50 mm), insulated with CSR Bradford sectional glass wool pipe insulation and fitted with a

Chapter 8: Basket Impeller Column 396 Digital RPM Stirrer

Rectifying Section Partial 9 T/s Condenser 8 & Collector Reactive 7 T/s Section 6

5 T/s 4 Stripping 3 T/s Section 2 Pump 1.

1 T/s

T/s

Reboiler Pump 2.

Figure 8.15 Multistage Basket Impeller Column Setup

1/4" shaft 1/8" inner rod d 30 NPT¼" BSP ¾" 4 mm, outer tube

150

5 5 25 10 %OA

dh 4 mm

5 %OA 4 mm id 50, dh 10 od 57

BSP ¾" a. b. c. d.

Figure 8.16 Multistage Basket Impeller Column Internals (a) Column Ends (b) Basket Cluster (c) Tray Support and Arrangement of One Stage (d) Tray Open Area

Chapter 8: Basket Impeller Column 397 water-cooled (at approximately 3 L min-1) glass condenser and a reboiler (a 5 L round bottom flask containing appropriate solution heated by an Electrothermal electric mantle, model MX).

Each stage consists of a polycarbonate tray perforated with 4 mm holes in an equilateral triangular pattern to the required open area and a 4-arm basket impeller made from 304 stainless steel mesh (315 aperture) installed on a rotating shaft running through the entire column. The shaft was powered by a Heidolph digital stirrer (model RZR 2021, 0 to 2000 rpm) mounted at the top end. Four tray supports consisting of an internal rod and sleeve tube of the same length as the tray spacing (150 mm) were equally spaced around the tray perimeter and also acted as baffles. The baskets were made rechargeable with the same bolts, which held them to the support ring. Since both tray and column were fabricated from polycarbonate, the maximum operating temperature in the basket impeller column was limited to 453 K. Table 8.05 summarises the parameter specifications for the rig.

Table 8.05 Physical Characteristics of the Basket Impeller Column Basket Impeller Column Details Value Overall Column Height, (m) 1.8 Internal Diameter, (mm) 50 Number of Trays 9 Tray Spacing, (mm) 150 Feed Tray (#1 bottom) 4 Tray Details % Open Area, %OA 5.8, 11.5 Hole Diameter, (mm) 4 Plate thickness, (mm) 2 Basket Details Number of arms in cluster 4 Arm height, (mm) 30 Arm width, (mm) 10 Arm length, (mm) 10 Cluster diameter, (mm) 32 Total Volume, (mL) 10 Packing Details

Amberlyst 15 dry dp , (μm) 363 Glass beads dp , (μm) 850

Chapter 8: Basket Impeller Column 398 8.3-2 Case Study A. System Characterisation

Experiments to determine the overall column efficiency, EOC, were performed using an ethanol-TBA mixture. The VLE data of Aucejo et al (1999) were employed in the calculation of the number of theoretical plates to achieve specified column exit composition. This system exhibits ideal VLE which are relatively flat requiring 45 equilibrium stages at total reflux for separation between xEtOH 0.05 and 0.95. This characterisation is useful in providing a qualitative guide to the region of high mass transfer rates for any subsequent catalytic distillation, irrespective of the specific reaction chemistry provided there is similarity in hydrodynamic and component properties, in this case alcohol dehydration. Distillation runs were carried out under total reflux following the procedure given by Subawalla et al (1997) with the reboiler initially charged with a 2L solution of 40 mol % ethanol (balance TBA). Two tray types with different percentage open areas (%OA) of 5.8 and 11.5 % were employed, the impeller speed, Ω, varying between 0 to 500 rpm. Boilup rate was adjusted to give vapour flow rates with F-factor over the range 0.1 to 0.5 Pa0.5. For these runs, the baskets were filled with 850 μm non-porous glass beads.

8.3-2-2 Fluid Dynamics under Separative Conditions

Froth height and pressure drops were measured manometrically. Figure 8.17 shows that increasing impeller speed initially appeared to promote froth formation because of better gas recirculation, bubble break-up and dispersion. Even so, this advantage disappeared around 100 rpm (depending on F-factor and tray type) such that thereafter froth height decreased with impeller speed probably because of increased liquid weeping through the holes as the basket spins on the tray forcing the clear liquid down through the tray holes. This phenomenon would not be apparent in a conventional dualflow column where frothing is due primarily to momentum exchange between the counter-currently flowing fluids.

For a dualflow tray, the froth height, hf, is related to the equivalent clear liquid height, hL, on the tray via:

hhLf=−(1 ε froth) (8.22) where εfroth is the gas fraction (porosity) in the froth.

Chapter 8: Basket Impeller Column 399 90 F = 0.100 80 F = 0.111 70 F = 0.119 60 F = 0.124 50 (mm)

f 40 h 30 20 10 0 0 100 200 300 400 500 Ω (rpm) (a)

120 F = 0.261 100 F = 0.324 F = 0.435 80

60 (mm) f h 40

20

0 0 100 200 300 400 500 (b) Ω (rpm)

Figure 8.17 Froth height, hf, variation with basket speed for different trays, (a) 5.8 %OA and (b) 11.5 %OA

Chapter 8: Basket Impeller Column 400 In a non-rotating system, Mahendru and Hackl (1979) proposed:

0.2 2 − ⎡U SGρ ⎤ ε froth =−1.0 0.0946 ⎢ ⎥ (8.23) ⎣ ghLLρ ⎦

and Xu et al (1994) have suggested:

b2 ⎡ LMn U ρρ/ 0.5 ⎤ ⎢()( SG ( air )) ⎥ hbL = 1 0.42 (8.24) ⎢ b3 ⎥ ρθLTh()td/ ⎣⎢ ⎦⎥

0.5 Since GL = LM and F=USρG Equation (8.24) becomes:

b ⎡ GFn ρ −0.5 2 ⎤ ⎢ ()Lair()⎥ hbL = 1 0.42 (8.25) ⎢ ρθb3 td/ ⎥ ⎣ LTh()⎦

-2 -1 where GL is the superficial liquid mass velocity (kg m s ). However, the empirical parameters, b1, b2, b3 and n appear to be dictated by the column dimensions. For example, while Xu et al (1994) found that b1 = 0.006, b2 = 1.75, b3 =1.90 and n = 0.3162 -0.25 θ ; while Garcia and Fair (2002) found that b1 = 0.01728, b2 = 1.0, b3 = 1.5 and n = 1.5 4.3θ . Combining Equation (8.22) and (8.23) gives the expression for hL in terms of hf as:

−0.25 2 1.25 ⎡ F ⎤ hhLf= 0.0524 ⎢ ⎥ (8.26) ⎣ gρL ⎦

since hf were measured values in this study, Equation (8.26) was used to estimate the corresponding equivalent clear liquid height for the stationary basket (Ω=0 rpm) in order for comparison with conventional dual flow trays as displayed in Table 8.06.

Chapter 8: Basket Impeller Column 401

Table 8.06 Comparison of hL Predicted by Equation (8.26) and Correlations of Xu et al (1994) and Garcia and Fair (2002).

-05 -04 %OA F hf (m) hL (m) hL ×10 hL ×10 (m) Xu (m) Garcia 0.058 0.100 0.040 0.028 1.146 1.889 0.058 0.111 0.050 0.034 1.487 2.114 0.058 0.119 0.058 0.039 1.755 2.278 0.058 0.124 0.076 0.052 1.968 2.396 0.115 0.261 0.042 0.020 1.432 1.675 0.115 0.324 0.048 0.021 5.690 2.163 0.115 0.435 0.075 0.028 11.170 3.051

Alternative estimates of hL from Equation (8.25) based on Xu et al’s (1994) and Garcia and Fair’s (2002) parameters were also obtained as displayed in column [4] and [5]. It is evident that hL increased with vapour velocity, F-factor, irrespective of the expression employed (cf Equation (8.25) or (8.26)). Nonetheless, as may be seen in Table 8.06, the values predicted by Equation (8.25) are considerably smaller than those from Equation

(8.26) suggesting that the b1 to b3 values used may not be applicable to the narrow (height : diameter ratio = 36) laboratory distillation column employed in this study.

Additionally, sensitivity analysis showed that hL was invariant with n, but more strongly influenced by b1 to b3. In order to estimate hL in a spinning system, the gas porosity in the froth, εfroth, was obtained from our previous correlation for stage gas holdup Equation (8.14) using the holdup definition of Equation (8.10):

0.986 0.337− 0.470 0.3272 ε froth = 0.00108ReGLIII Re Re()hw / for Ω≠0 (8.27)

which combines with Equation (8.22) to yield hL values. Thus, a more general version of Equation (8.26) to account for the variation of hL in Ω was proposed as:

⎡⎤n −0.5 b2 GFLair()ρ hb=Ω+Ω+⎢⎥⎡ a2 a 1⎤ (8.28) L 112b 0.42 ⎣ ⎦ ⎢⎥ρθ3 td/ ⎣⎦LTh()

The data for both stationary and rotating baskets were used to obtain b1, b2, b3, a1 and a2 via non-linear regression using POLYMATH 5 as b1 = 2.08961, b2 = 1.137, b3 = 1.736, -06 -03 a1 = -6.23 × 10 and a2 = 2.53 × 10 . Values of n were obtained from Garcia and Fair

Chapter 8: Basket Impeller Column 402 (2002). Equation (8.28) gave a correlation coefficient of 0.82. Figure 8.18, the parity plot, shows the good agreement between observed and predicted hf based upon Equation (8.27) and (8.28). Clearly, the expression reduces to the form of Equation (8.25) for non-rotating basket operation (Ω=0). In particular, it predicts a maximum in hL with impeller speed, however, because it is not mechanistically-based, we have not assigned physical meaning to the empirical parameters, b1, b2, b3, a1 and a2.

Measured column pressure drop plotted as a function of the basket impeller speed is shown in Figure 8.19 for the two tray types. The decrease in pressure drop in a rotating system could be due to cycloning, where the liquid is pushed outward by centrifugal forces forming a film on the column wall, while the gas is promoted to flow upwards at the centre. However, in a baffled wall column such as depicted in Figure 8.16 (c) the tray supports, which acted as baffles would prevent the formation of a permanent downward flowing film on the column wall and hence minimise the cycloning effect. In fact, visual inspection during the runs confirmed an even distribution of froth above each perforated tray. Thus, the decrease in pressure drop is primarily due to a reduction in equivalent clear liquid height, hL, defined in Equation (8.28). Consistent with the data on froth height variation, these curves also show that column pressure drop decreased with increased basket impeller speed, before levelling out at about 200 rpm, where the clear liquid height on the tray is probably at a minimum because of weeping and dumping through the holes from the 2 trays (top and bottom) of the stage.

0.1

0.08

0.06 predicted f 0.04 h

0.02

0 0 0.02 0.04 0.06 0.08 0.1

hf observed

Figure 8.18 Parity plot between observed and predicted hf values

Chapter 8: Basket Impeller Column 403 4.0 F = 0.111 3.5 F = 0.119 3.0 F = 0.124

2.5

2.0 P (kPa)

Δ 1.5

1.0

0.5

0.0 0 100 200 300 400 500 Ω (rpm) (a)

6 F = 0.261 5 F = 0.324 F = 0.435 4

3 P (kPa) Δ 2

1

0 0 100 200 300 400 500 Ω (rpm) (b)

Figure 8.19 Influence of basket impeller speed on column pressure drop , ΔP (a) 5.8 %OA and (b) 11.5 %OA

Chapter 8: Basket Impeller Column 404 8.3-2-3 Overall Column Efficiency

For the non-reactive distillation of the ethanol-TBA mixture, aliquots were taken from stages 1 and 9 every 30 minutes and analysed until repeated sampling indicated that steady state had been attained (typically after about 3 hours). The average column efficiency, EOC, to achieve the observed separation was obtained from:

theoretical number of plates E = (8.29) OC actual number of trays = 9

The theoretical number of trays was obtained via a numerical McCabe-Thiele method. This was done by stepping upwards and downwards off the VLE curve using a 45o operating line and starting from the experimentally measured compositions at the appropriate end (top or bottom) until the concentration at the other end had been reached. The average number of trays in both directions was then accepted as the theoretical number of stages to effect the observed separation.

Figure 8.20 reveals that the overall column efficiency, EOC, is a function of the impeller speed for both tray types at all the vapour velocities examined. Compared to a non- rotating system, EOC value initially increased as impeller speed increased indicating better gas-liquid mass transfer coefficient with mechanical agitation. However, beyond the optimum speed, located at around 200 rpm, separation efficiency dropped because of intense liquid weeping through the tray holes at these high speeds and a rapid disruption of the vapour jet formation, which combine to give a reduced froth zone hence poorer gas-liquid contacting. Incidentally, this behaviour is parallel to the pattern seen in hf – an indication that separation efficiency may be foreshadowed by the hydrodynamic behaviour. Even so, the data obtained under a stationary impeller (Ω = 0) suggest that EOC increased with vapour velocity, F-factor, in agreement with the findings of Xu et al (1994) and Garcia and Fair (2002). Moreover, this investigation showed that EOC values close to 100% are possible even with the relatively low F-factor employed because of improved interphase mixing and mass transfer in the spinning basket dualflow tray column.

Chapter 8: Basket Impeller Column 405 1.4 F = 0.100 1.3 F = 0.111 F = 0.119 1.2 F = 0.124 1.1 OC E 1.0

0.9

0.8

0.7 0 100 200 300 400 500 Ω (rpm) (a)

1.4 1.3 F = 0.261 1.2 F = 0.324 1.1 F = 0.435 1.0

OC 0.9 E 0.8 0.7 0.6 0.5 0.4 0 100 200 300 400 500 Ω (rpm) (b)

Figure 8.20 Overall Column Efficiency, EOC, as a function of stirrer speed: (a) 5.8 %OA and (b) 11.5 %OA

Chapter 8: Basket Impeller Column 406 In conventional dualflow tray distillation, much higher F-factor values (>1.0) are required to accomplish EOC estimates close to unity Garcia and Fair (2002), and this would require a much higher reboiler duty and lead to a greater propensity to flooding. A quadratic expression offers the simplest representation over the range of vapour flow rates examined (0.1 ≤ F ≤ 0.5), thus:

2 EOC =+Ω+Ωα01αα 2 (8.30) where,

−03 αθ0 =−1.08 2.51 −× 7.97 10 F (8.31)

−−06 06 −05 αθ1 =−3.36 × 10 + 6.14 × 10 − 6.17 × 10 F (8.32)

−−03 02 −02 αθ2 =×2.08 10 −× 4.69 10 +× 1.17 10 F (8.33)

and a regression coefficient of 0.9635 with an optimum basket speed, Ωopt = -α2/2α1. Thus, for a catalytic distillation system it would be desirable to operate under conditions, which maximise EOC in order to ensure that separation-limited effects do not affect reaction rate or product selectivity.

Chapter 8: Basket Impeller Column 407 8.3-3 Case Study B. Application to Catalytic Distillation Catalytic distillation study was carried out using the dehydration of TBA. The basket impellers of stages 5 to 8 constituting the reactive zone were filled with Amberlyst 15 catalyst particles (average dp=363 μm). Runs were conducted at 200 and 500 rpm while feed rate was varied via a Masterflex Quickload peristaltic pump (Pump 1). The feed was an equimolar mixture of TBA and water – a composition to the right of the distillation boundary of the conventional residue curve generated using the UNIFAC thermodynamic model. The boilup rate was maintained at 10-2 mol s-1 with a reboiler duty of 1080 watts. The reboiler residue was withdrawn using Pump 2. Liquid samples were collected from every second tray for TBA and water analysis. However, a separate set of runs was also carried out in a 500 mL stirred slurry reactor to determine the intrinsic kinetics of TBA dehydration at 355 ± 1.5 K. This was used for base comparison and to evaluate the performance of the catalytic distillation column.

8.3-3-1 TBA Dehydration Kinetics

In order to evaluate the performance of the basket impeller column for the catalytic dehydration of TBA under distillation conditions, it was necessary to obtain the kinetics. The kinetics were determined at 355 K under boiling conditions with a similar experimental approach as adopted by Frilette (1964). The experiments were carried out in a 500 mL stainless steel mechanically agitated slurry reactor. Figure 8.21 shows the effect of stirring rate on the rate of TBA dehydration and it is apparent that external mass transfer resistance was negligible at impeller speed greater than 100 rpm. The non- zero intercept at Ω = 0 rpm arose from the stirring occasioned by the mild boiling conditions of the liquid phase mixture at the reaction temperature (T=355K).

Earlier studies have shown that internal mass transport limitation was non-existent with particle size less than 100 μm, Honkela et al (2004) and Gates et al (1973). However, there is evidence for significant particle swelling in an aqueous TBA mixture, Heath and Gates (1972) (typical of what will be obtained on a dualflow distillation column). Hence, it was necessary to establish a functional relationship for the effectiveness factor in terms of the particle size (anhydrous state) and water concentration in the reactor mixture.

Chapter 8: Basket Impeller Column 408 6.0

5.9

5.8 ) -1

.s 5.7 -1 5.6

(mol.g 5.5 +05 5.4 .10 5.3 dehyd r 5.2

5.1

5.0 0 100 200 300 400 500 600 Ω (rpm)

Figure 8.21 Effect of stirring speed on TBA dehydration rate over Amberlyst-15

1.2 i Water

) 1.0

-1 i TBA .s -1 0.8

(mol.g 0.6 +04 10

. 0.4 dehyd r 0.2

0.0 0 5 10 15 20 25 -1 Ci (mol.L )

Figure 8.22 Kinetics of TBA dehydration rate

Chapter 8: Basket Impeller Column 409 Rate measurements at 2 different dp values (75 and 363 μm) were obtained at various water concentrations and Equation (4.54) of Section 4.3-2 was numerically solved for the data represented in Table 8.07. The inhibition effect of water on η provides additional impetus for the application of catalytic distillation to this type of reaction system.

Table 8.07 The Effectiveness Factor with CW and dp

-1 -05 - dp (μm) CW (mol.L ) r ×10 (mol.g η 1.s-1) 75 0.57 10.496 0.994 75 2.56 8.189 0.992 75 3.72 7.670 0.982 363 0.57 9.363 0.887 363 2.56 7.042 0.853 363 3.72 5.713 0.732

Further experiments were carried out with 75 μm particles using an agitation speed of 200 rpm to determine the effect of TBA and water concentration on reaction rate. As may be seen from Figure 8.22 while rate increased almost linearly with TBA concentration, water has a detrimental effect on the kinetics. This behaviour has been attributed to the competitive chemisorption of the TBA and water on anhydrous, non- ionised sulphonic groups, Frilette (1964). Abella et al (1999) have also observed that reaction rates were different over dry and wet resins because of particle swelling in the presence of water. They proposed a rate equation, which incorporated the inhibition of water without TBA chemisorption namely:

kC1 TBA −=rTBA (8.34) 1+ KCWW

Kato et al (1996) have also reported that TBA dehydration in a semi-batch reactor may be represented by:

kC1 TBA −=rTBA 2 (8.35) 1++KCIB IB KC W W

to account for isobutylene adsorption on acid sites.

Chapter 8: Basket Impeller Column 410

Gates and Rodriguez (1973) postulated that the reaction proceeds via a carbonium ion mechanism at low sulphonic (-SO3H) concentration but at high concentrations a concerted mechanism with TBA hydrogen bridged in a network of –SO3H groups was

Chapter 8: Basket Impeller Column 411 Table 8.08 Summary of Recent Kinetic Models for the Dehydration of TBA

+06 2 Source LH expression for Mechanistic Features k1 ×10 KTBA KW R of (-rTBA) model Abella kC Only water is adsorbed TBA 9.749 - 5.082 0.9823 1 TBA et al 1+ KC reacts from the liquid phase. WW

Kato et kC Water and IB are absorbed but 8.537 - 0.185 0.7445 1 TBA al 1++KC KC2 two molecules of water adsorb IB IB W W on a single site. TBA is not chemisorbed.

Gates kC TBA adsorption with two 2.884 -0.067 0.014 0.9335 1 TBA et al 1++KC KC2 molecules of water TBA TBA W W chemisorbed.

This Competitive adsorption 2060 1.313 53.05 0.9741 kC1 TBA Study 2 between water and TBA. Rate 1++KC KC ()TBA TBA W W controlling step is the vacant site-assisted dehydration of chemisorbed TBA.

Chapter 8: Basket Impeller Column 412 implicated. Given the essential features of the reaction mechanism found in the literature, the following sequence of steps may be proposed:

TBA+= S TBA. S () molecular adsorption of TBA (M.1a)

or TBA+= S IB. S + W ( single-site dissociative TBA adsorption) (M.1b)

TBASSWSIBS.+= . + . ( vacant site-assisted dehydration) (M.2a)

or TBA. S=+ IB W . S () dehydration of chemisorbed TBA (M.2b)

WS.=+ W S () desorption of water (M.3)

IB. S=+ IB S () desorption of isobutylene (M.4)

where S is a protonated active site ( H+ ). Previous studies, Gates et al (1973), indicate that water is more strongly adsorbed than isobutylene on ion-exchange resin catalysts because of its relatively higher polarity, hence only water and TBA would be considered the most abundant surface intermediates. Formal Langmuir-Hinshelwood treatment of Equations (M1) to (M4) yields:

kC1 TBA −=rTBA 2 (8.36) ()1++KCTBA TBA KC W W

assuming that the rate-determining step is the irreversible vacant site-assisted surface dehydration of chemisorbed tert-butyl alcohol (cf. Equation (M.2a)) with rapid equilibrium in steps (M.1a) and (M3). In fact, earlier kinetic models shown in Table 8.08 could be readily derived from the proposed mechanism depending on the

Chapter 8: Basket Impeller Column 413 assumptions made. Although the kinetic model by Abella et al (1999) gave a high correlation coefficient, the requirement that TBA reacts from the fluid phase is inconsistent with the strong evidence for TBA adsorption prior to dehydration seen in several studies and hence incompatible with a Langmuir-Hinshelwood rate law formulation. On the other hand, the model of Gates et al (1973) gave a negative equilibrium adsorption constant for TBA – a physical oddity. Thus, Equation (8.36) was accepted as the best model in view of the rationale behind its assumptions and the associated high correlation coefficient (0.974).

8.3-3-2 Catalytic Distillation Runs

Catalytic distillation runs were then conducted using the TBA-water feed mixture (50 mol % TBA) supplied to tray 4 at 353 K. The reboiler was maintained at 375 K and tray 9 (top tray) was kept at 313 K by the overhead total reflux stream. Baskets in trays 5 to 8 were filled with 363 μm Amberlyst 15 catalyst particles while baskets in the non- reactive section contained glass beads. Initial runs revealed that conversion increased with decreased flow rate. A similar behaviour had been reported by Abella et al (1999). In order to achieve complete conversion we found that TBA feed rate had to be varied between 0.73 to 2.50 mmol s-1 for the 200 and 500 rpm runs. The basket impeller column would operate within the limits of a well-stirred reactor and plug flow reactor. Hence, an estimate of the expected conversion from either of these ideal reactors for the experimental feed rate (0.75 mmol s-1), would be instructive in the characterisation of the reaction behaviour in a reactive-separation system. Simple application of the ideal CSTR model based on the slurry kinetics of Equation (8.36) yields:

1− X = 3.544 (8.37) −−02 02 2 XXX⎣⎦⎡⎤1+× 0.531 10() 1 ++× 1.313 10 () 1 −

with the feasible root at XCSTR = 0.20. Similar treatment for the PFR provides:

Chapter 8: Basket Impeller Column 414 ⎧⎫ ⎪⎪ ⎪⎪515.040⎣⎦⎡⎤()() 1−− X 2ln 1 −+ X 637.381X −− 1 ⎪⎪ 22 ⎪⎪4 ⎛⎞⎛⎞XX5 ⎨⎬4.854×−+++−=× 10 ln() 1 X 6.763⎜⎟⎜⎟ X 8.37 X 1.37 10 (8.38) ⎪⎪⎝⎠⎝⎠22 ⎪⎪2 ⎡⎤()1X− ⎪⎪+−−−−−1.366⎢⎥ 4()() 1 X 4ln 1 X 3.5 ⎪⎪2 ⎩⎭⎣⎦⎢⎥ which gives XPFR = 0.35. Thus, the new catalytic distillation dualflow tray BIC reactor is better than either of the two ideal reactors, hence, the promotive effect of the combined reaction and distillation in a single vessel.

Figure 8.23 shows that the liquid phase concentration profiles are approximately linear with respect to axial position (tray number) suggesting that the overall reactive distillation performance may be regarded as the arithmetic sum of contribution from each tray in stages 5 to 8. A conservative performance indicator of the new catalytic distillation reactor the enhancement factor, ξ, defined:

rBIC,obs ξ = 8 (8.39) ∑ rslurry, i i=5

where rBIC,obs is the observed overall TBA reaction rate in the new basket impeller th column and rslurry, i is the effective reaction rate on the i tray at the corresponding TBA concentration based on the previously obtained slurry reactor kinetics in Equation (8.38). As a result,

rrslurry kinetics =−η ( TBA ) (8.40) where the multiplication by the effectiveness factor, η, arose to account for the water- induced swelling of the catalyst particle and possible mass transfer effect variation from tray-to-tray as the water concentration changes. From the data in Table 8.07, η correlates with particle size, dp, and water concentration, CW, as:

Chapter 8: Basket Impeller Column 415 ⎛⎞⎛⎞1.04 7.6 −04 η =−⎜⎟⎜⎟0.012 −−× 2.313 10 d p for CW>0 (8.41) ⎝⎠⎝⎠CCWW

This expression was combined with Equation (8.40) to evaluate ξ in Equation (8.39). The plot of ξ versus TBA feed flow rate shown in Figure 8.24 suggests an optimum enhancement factor of about 1.3 was achieved within the range of flow rate examined. By comparison, a non-rotating system with same reaction rate law and operated under identical hydrodynamics gave an ξ value of 0.74 (using data from Abella et al (1999).

10

9

8 xW 500 rpm xTBA 500 rpm 7 xW 200 rpm 6 xTBA 200 rpm

5

plate no. 4

3

2

1

0 00.51

xi

Figure 8.23 Concentration profile obtained at Ω=200 and 500 rpm for TBA feed rate of 10-03 mol s-1

Chapter 8: Basket Impeller Column 416 This translates to a 75% improvement for the new basket impeller catalytic distillation column. The improvement in the spinning basket impeller is because of superior separation efficiency of the reaction species as signalled by the range of Damkohler number, Da, found in this study (<1.0). Specifically, the Da is a measure of the relative rate of reaction to the transport (convective or diffusive) rate.

The modified Damköhler number for catalytic distillation within a basket impeller column was defined as:

8 ρIafAh∑ rslurry kinetics i Da = i=5 (8.42) V where V is the vapour flow rate through the dualflow tray. As may be seen from Figure 8.25, over the range of liquid flow rates examined, the transport rate was higher than the reaction rate in all cases since Da values were less than unity. This suggests that the reaction was not limited by phase equilibrium. In particular, the thermochemical equilibrium constant for the reaction is also low (K = 0.04 mol L-1 at the reaction temperature). It is therefore evident that the new spinning basket catalytic distillation reactor permits application to systems with low Da and low chemical equilibrium constant parameters which would be a disadvantage in the conventional fixed bed catalytic distillation column(7-8). Furthermore, it is interesting that the best conditions for separation efficiency and favourable catalytic distillation enhancement factor were found with a basket speed of 200 rpm and similar hydrodynamics (0.1

Chapter 8: Basket Impeller Column 417 1.4

1.2

1.0

0.8 ξ 0.6

0.4 200 RPM 0.2 500 RPM 0.0 0.0 1.0 2.0 3.0

03 -1 x10 FeedTBA (mol.s )

Figure 8.24 Enhancement factor at various TBA feed rates

0.6

200 RPM 0.5 500 RPM

0.4

0.3 Da

0.2

0.1

0 0123 03 -1 x10 FeedTBA (mol.s )

Figure 8.25 Damköhler number as a function of TBA feed rates

Chapter 8: Basket Impeller Column 418 8.4 Concluding Remarks and Further Considerations

8.4-1 Conclusions For the combination of rotating basket reactor and dualflow tray it was found that it was possible to obtain tray loading and buildup of froth for flowrates typical of distillation practice. Liquid and gas holdups and froth height eventually decreased as Ω (rpm) was increased, due to an increase in the rate of weeping caused by stronger forced circulation. Empirical correlations were proposed for liquid holdup, gas froth holdup, froth height, and pressure drop. The power used to stir the froth under aerated conditions was measured and found to increase with Ω (rpm) as may be expected. A suitable correlation was proposed for power consumption.

A multistage version of the Basket Impeller Column (BIC) was fabricated and it was demonstrated that distillative separation and reactive separation are feasible within this catalytic distillation reactor design. The overall column efficiency at relatively low column loading could be increased to values close to unity through the variation of speed of rotation. The speed of rotation could also be used to directly vary Damköhler number under catalytic distillation conditions. Thus, the BIC design has potential as a catalytic distillation application development tool. Using a dh = 4mm and 10 %OA the speed of 200 rpm is preferred for both conventional and catalytic distillation.

8.4-2 Potential Applications of the BIC For the purpose of synthesis of TBA the Basket Impeller Column can be applied as a either a countercurrent reactor in the or as a catalytic distillation. Applied as countercurrent reactor hot water could be fed at the top and a C4 stream containing isobutylene at the bottom. The BIC should allow better holdup and longer liquid residence times than conventional pack fixed bed reactors while obtaining near complete mixing and liquid/catalyst renewal. In the CFBR study it could be seen that for Bale packing mixing was strongly a function of liquid flowrate below the loading point. Further, TBA reaction rates obtained were also a strong function of mass transfer limitations governed by localised mixing and hence it was difficult to tailor holdup to the equilibrium requirements of the hydration. Operated as a countercurrent reactor the

Chapter 8: Basket Impeller Column 419 BIC will also have a reactive separation function as the other C4 components such as n- butene will be separated from isobutylene through reaction.

The BIC can be applied for the hydration of isobutylene under pressurised conditions. The strong mixing offered by the rotating basket impellers could disperse the separate liquid phases, which are likely to form. This inter-mixing would improve mass transfer between the phases formed. In order to gain miscibility within the system of hydration of isoamylene, Gonzalez and Fair (1997) used acetone mole fractions of typically x = 0.9, thus flooding the system with solvent, which in turn led to adol condensation of acetone. If the focus on the use of solvent is placed upon activity alteration of TBA and water and not solubility of isobutylene then lower amounts of solvent should be required.

For the hydration of isobutylene the magnitude of reaction rates expected should be of a similar order to those determined in the basket slurry reactor employed in the kinetics study of Chapter 4. This reactor could be considered to be a single stage of the BIC with infinite residence time in the liquid phase. For the isobutylene hydration system it is important to attain an overall effectiveness factor as close to unity as possible as with reaction rates of the order of magnitude of 1 × 10-07 to 1 × 10-06 mol/g.s the reaction is relatively slow.

The potential benefits of the BIC may also find application in other systems struggling to gain the complete benefits of catalytic distillation due to the mass transfer or separation limitations of current static hardware designs. Such systems include adol condensation of acetone, butanol and acetic acid esterification to butyl acetate and fatty acid esterification.

In the case of adol condensation of acetone to diacetone alcohol the system is plagued by liquid-liquid mass transfer limitations and there are reaction selectivity issues arising from the consecutive reaction of diacetone alcohol to mesityl oxide, Nicol (2003). Further, the inhibition effect of water is very strong and good separation is required within the reactive zone. Improved mixing and better separation within the reaction zone should assist the removal of water. The strongly basic ion exchange resins in the

Chapter 8: Basket Impeller Column 420 OH- form, as required for this reaction, such as Amberlyst A26, are particularly susceptible to thermal degradation. The stronger the basic nature of the catalyst the more likely it is to deactivate. Thus, it is very important to maintain a uniform temperature distribution avoiding hot spots typically associated with liquid maldistribution in conventional packing. The good mixing of the BIC may allow the use of Amberlyst A26 (maximum temperature limit of 333 K) over lower strength resins such Amberlyst A21 or Amberlite IRA 900.

The benefits for the butyl acetate system may include the prevention of the side reaction to dibutyl ether and water through better separation within the actual reaction zone as opposed to larger non-reactive stripping and rectifying sections as proposed by Gangadwala et al (2004). The complication of this side reaction has only come to light recently and the toxicity of the product makes its prevention crucial if the butyl acetate is to be used as a common solvent.

For fatty acid esterification the benefits of employing the BIC would arise through better mass transfer in the case of these large high viscosity molecules. Further, the stirred system may encourage better water removal from the two liquid phase mixtures formed. The proposed systems employing azeotropic entrainers such as that of Dimian et al (2003) for the esterification of propanol and lauric acid may benefit through better control of holdup.

8.5 Nomenclature

%OA Percentage Open Area, dimensionless a0, a1 Constants as defined 2 Aa Column cross sectional area, (m ) 2 Ah Total open area of tray, (m ) b1, b2 Constants as defined -1 CAS Surface concentration of component A, (mol.L ) -1 Ci Concentration of component i, (mol.L )

CWP Weise-Prater criterion d Column diameter (m) dh Hole diameter, (m)

Chapter 8: Basket Impeller Column 421 dI Impeller diameter, (m) 2 De Effective diffusivity, (m .s) Da Damköhler number, Equation (8.42)

EOC Overall column separation efficiency 0.5 F Capacity factor, (US.ρ ) 3 FG Volumetric Gas Flowrate (m /s) G Gas Superficial Mass Velocity, (kg/m2.s) -2 -1 GL liquid mass flux, (kg.m .s ) g Acceleration because of gravity, (m.s-2) hFroth Froth height, (mm) hI Impeller height, (m) hL Equivalent clear liquid height, (m)

HL Liquid phase holdup, volume fraction, dimensionless

HG Total Gas phase holdup, volume fraction, dimensionless

HG, Froth Froth Gas phase holdup, volume fraction, dimensionless

HG, DH Disengaging Height Gas phase holdup, volume fraction, dimensionless k Rate constant, (L.g-1.s-1)

Ki Equilibrium adsorption constant of component i L Liquid molar flux, (kmol.m-2.s-1) M Molecular mass, (kg.kmol-1) n Exponent equations

PG Power (W) for Gasified Liquid.

P0 Power (W) for Flat Liquid ΔP Pressure Drop, (Pa) Q Volumetric Flowrate (m3/s)

ReG,L Sieve Tray Reynolds number, [ρUdh/μ]

ReI Impeller Reynolds number, [ρ(rps)(wI*hI)/μ] S Active site tT Plate thickness, (m) T Temperature, (K) U Sieve Plate Orifice Velocity, [Q/((πd2/4) × %OA)], (m/s) V Volume of Liquid (m3) wI Impeller width, (m)

Chapter 8: Basket Impeller Column 422 X Conversion

Greek Symbols

α0, 1,2 Constants in the empirical model for EOC

εfroth Froth porosity, (vol/vol) φ Thiele modulus η Effectiveness factor

θ Ratio of areas, (Ah/Aa) ρ Density, (kg.m-3) -3 ρI Impeller packing density, (kg.m ) ξ Enhancement factor Ω Rotational speed (rpm)

8.6 Literature Cited Abella, L. C., P. A. D. Gaspillo, M. Maeda and S. Goto “Kinetic Study on the Dehydration of tert-Butyl Alcohol Catalyzed by Ion Exchange Resins.” Int. J. of Chem. Kinetics 31(12): 854-859. (1999). Abella, L. C., P. D. Gaspillo, H. Itoh and S. Goto “Dehydration of ter-butyl alcohol in reactive distillation.” J. Chem. Eng. of Japan 32(6): 742-746. (1999). Akagi, Y., K. Okada, K. Kosaka and T. Takahashi “Liquid weeping accompanied by bubble formation at submerged orifices.” Ind. Eng. Chem. Res. 26: 1546-1550. (1987). Aucejo, A., S. Loras, R. Munoz and L. M. Ordonez “Isobaric vapour liquid equilibrium for binary and ternary mixtures of ethanol + 2-methyl-2-propanol and 2- methylpentane + ethanol + 2-methyl-2-propanol.” Fluid Phase Equilibria 162: 2411. (1999). Billet, R. “Separation Tray without Downcomers.” Chem. Eng. Tech. 24(11): 1103- 1112. (2001). Bruijn, W., K. v. t. Riet and J. M. Smith “Power Consumption with Aerated Rushton Turbines.” Trans. IChE 52: 88-104. (1974).

Chapter 8: Basket Impeller Column 423 Carberry, J. J. "Chemical and Catalytic Reaction Engineering". NY, McGraw-Hill Book Company. (1976). Colburn, A. P. “Effect of entrainment on plate efficiency in distillation.” Ind. Eng. Chem. 28(5): 526-530. (1936). Dimian, A., F. Omota and A. Bliek “Entrainer enhanced reative distillation.” Chemical Engineeering and Processing in press. (2003). Fair, J. R. (1987). Distillation. Handbook of Separation Process Technology. R. W. Rousseau. Brisbane, John Wiley & Sons: 229-239. Frilette, V. J., E. B. Mower and M. K. Rubin “Kinetics of dehydration of tert-butyl alcohol catalysed by ion exchange resins.” J. Catalysis 3(1): 25-31. (1964). Furzer, I. A. “Froth height on dualflow trays with ternary azeotropic system of ethyl acetate-ethanol-water.” Ind. Eng. Chem. Res. 39(5): 1430-1436. (2000). Furzer, I. A. “Froth heights on dualflow trays with heterogeneous binary azeotropic system and a heterogeneous ternary system with a homogeneous azeotrope.” Ind. Eng. Chem. Res. 40(22): 4951-4966. (2001). Gangadwala, J., A. Kienle, E. Stein and S. Mahajani “Production of Butyl Acetateby Catalytic Distillation:Process Design Studies.” Ind. Eng. Chem. Res. 43: 136- 143. (2004). Garcia, J. A. and J. R. Fair “Distillation Sieve Trays without downcomers:Prediction of performance characteristics.” Ind. Eng. Chem. Res. 41: 1632-1640. (2002). Gates, B. C. and W. Rodriguez “General and Specific Acid Catalysis in Sulfonic Acid resin.” J. Catalysis 31: 27-31. (1973). Gonzales, J. C., H. Subaealla and J. R. Fair “Preparation of tert-Amyl Alcohol in a Reactive Distillation Column. 2. Experimental Demonstration and Simulation of Column Characteristics.” Ind. Eng. Chem. Res. 36: 3845-3853. (1997). Gotze, L., O. Bailer, P. Moritz and C. V. Scala “Recative distillation with KATAPAK.” Catalysis Today 69: 201-208. (2001). Gupta, V. P. and J. M. Douglas “Diffusion and Chemical Reaction in Isobutylene Hydration within Cation Exchange Resin.” AIChE Journal 13(5): 883-889. (1967). Heath, H. W. and B. C. Gates “Mass Transport and Reaction in Sulfonic Acid Resin Catalyst: the Dehydration of t-Butyl Alcohol.” AIChE Journal 18(2): 321-326. (1972).

Chapter 8: Basket Impeller Column 424 Honkela, M. L., T. Ouni, A. Krause and I. Outi “Thermodynamics and kinetics of dehydration of tert-buty alcohol.” Ind. Eng. Chem. Res. 43(15): 4060. (2004). Ihm, S. K., M. J. Chung and K. Y. Park “Activity Difference between the Internal and External Groups of Macroreticular Ion Exchange Resin Catalysts in Isobutylene Hydration.” Ind. Eng. Chem. Res.(27): 41-45. (1988). Kato, Y., Y. T. Honda and A. Kanzawa “Kinetic measurement on the isobutene/water/tert-butyl alcohol heat pump: dehydration of tert-butyl alcohol.” Int. J. of Energy Research 20: 681-692. (1996). Kelleher, T. and J. R. Fair “Distillation studies in a high-gravity contactor.” Ind. Eng. Chem. Res. 35: 4646-4655. (1996). Krishna, R. “Reactive separations: More ways to skin a cat.” Chem. Eng. Sci. 57: 1491- 1504. (2002). Mahendru, H. L. and A. Hackl “Contribution to the design of sieve trays withou downcomers.” Inst. Chem. Eng. Symp. Ser. 56(3.2): 35-47. (1979). Miyahara, T., M. Kurihara, M. Asoda and T. Takahashi “Gas-liquid interfacial area and liquid-phase mass transfer coefficent in sieve columns without downcomer operating at high gas velocities.” J. Chem. Eng. Japan 23(3): 280-285. (1990). Myers, H. S. “A Versatile Fracxtionating Column.” Ind. Eng. Chem. 50(11): 1671- 1674. (1958). Nicol, W. “Comparing catalytic distillation to separate reaction and distillation for the production of diacetone alcohol.” Trans. I Chem. E. 81(Part A): 1026-1032. (2003). Peng, W. L., G. Yang and L. S. Fan “Experimental Studies of liquid weeping and bubbling phenomena at submerged orifices.” Ind. Eng. Chem. Res. 41: 1666- 1677. (2002). Quitain, A., H. Itoh and S. Goto “Reactive Distillation for Synthesizing Ethyl tert-Butyl Ether from Bioethanol.” J. Chem. Eng. of Japan 32(3): 280-287. (1999). Safinski, T. and A. A. Adesina “Development of a Novel Basket Impeller Dualflow Tray Catalytic Distillation Reactor.” Ind. Eng. Chem. Res. 44, 6212-6221 (2005). Safinski, T. and A. A. Adesina “Parametric Study of a novel basket impeller column.” Can. J. Chem. Eng. 81: 574. (2003).

Chapter 8: Basket Impeller Column 425 Sinnott, R. K. "Coulson and Richardson's Chemical Engineering". Sydney, Butterworth and Heinemann. (1997). Subawalla, H., J. C. Gonzalez, A. F. Seibert and F. J. R. “Capacity and Efficiency of Reactive Distillation Bale Packing: Modeling and Experimental Validation.” Ind. Eng. Chem. Res. 36: 3821-3832. (1997). Takahashi, T., T. Miyahara, T. Tano and Y. Akagi “Hydrodynamic characterisation of sieve trays having large free areas without downcomers.” J. Chem. Eng. 19(4). (1986). Tuchlenski, A., A. Beckmann, D. Reusch, R. Dussel, U. Weidlich and R. Janowsky “Reactive distillation industrial applications, process design and scale-up.” Chem. Eng. Sci. 56: 387-394. (2001). Wright, A. J. and D. L. Pyle “An investigation into the use of spinning cone column for in situ ethanol removal from yeast broth.” Process Biochemistry 31(7): 651-658. (1996). Xu, Z. P., A. Afcan and K. T. Chuang “Efficency of dualflow trays in Distillation.” Can. J. Chem. Eng. 72(August): 607-613. (1994). Yanagi, T. “Inside a Tray Distillation Column.” Chemical Engineeering Nov.: 120-129. (1990). Zhang, W. and R. B. H. Tan “The influence of liquid head and liquid circulation on weeping rates at a submerged orifice.” Chem. Eng. Tech. 26(11): 1169-1175. (2003).

Chapter 8: Basket Impeller Column 426 9. Conclusion and Recommendation

As part of this study the solubility of isobutylene, as a function of temperature and concentration within the water/TBA/ethylene glycol system, was determined. The solubility of isobutylene is ten times higher in TBA than in water and is twice as high in ethylene glycol than in water. Solubilities in water and ethylene glycol mixtures are approximately equal to the mole fraction based contribution of the pure solubility in each. The solubility within mixtures containing TBA is highly nonlinear and skewed towards that in pure TBA. The O’Connell-Prausnitz model for a two-way interaction mechanism adequately represents the solubility of isobutylene within ternary mixtures of water, ethylene glycol and TBA:

lnHxHxHxH=+ ln ln +ln + IB... mix W IB W EG IB EG TBA TBA (4.43) ααW.. EGxx W EG++ W TBA xx W TBA α EG . TBA x EG x TBA where:

−03 αWEG. =×3.026 10T − 14.14 (4.44)

−03 αWTBA. =−7.281 × 10T − 14.88 (4.45)

−02 α EG. TBA =−2.069 × 10T + 3.53 (4.46)

The use of this correlation is recommend for systems under atmospheric conditions.

The kinetics of isobutylene hydration, in the presence of the solvent ethylene glycol at the conditions of the proposed catalytic distillation column, were determined utilising a stirred basket reactor. The following heterogeneous kinetic model expression based upon the LHHW formalism, and an ER mechanism was proposed for the hydration of isobutylene in the presence of ethylene glycol:

Chapter 9: Conclusion and Recommendation 427 ⎛⎞−E kCC0 exp⎜⎟IB W ⎝⎠RT rTBA = (4.76) ⎛⎞⎛⎞a1 ⎛⎞a3 ⎜⎟1exp++++⎜⎟aC24WE exp⎜⎟ aCG ⎝⎠⎝⎠TT⎝⎠ where 2 -1 -1 -1 k0 = 345.65 (L .mol .g .s ) E = 19.5886 (kJ.mol-1) a1 = 4582.6 a2 = -6.205 a2 = 4662.2 a3 = -5.0915

This model is recommended for the evaluation of maximum, attainable reaction rates for the conditions of insignificant TBA, MET, and DET concentrations. It is recommended that CIB required, be estimated using Equation (4.43) to (4.46).

The solvent ethylene glycol is not inert, and MET is formed as by-product. The relative rate of TBA to MET formation is described by:

0.89 ⎛⎞−1127 ⎛⎞CW SRTBA=1.0exp⎜⎟⎜⎟ C W (4.67) ⎝⎠TC⎝⎠EG

Higher temperatures and lower solvent concentration favour selectivity towards TBA.

The solvent ethylene glycol is particularly effective in breaking the water/TBA azeotrope. Extractive distillation is feasible over the small scale Bale packing prepared specifically for this study. The molar feed rates of azeotrope and solvent, which avoid excessive reverse entrainment of TBA and entrainment of solvent, place the distillative system within the pre-loading regime. At zero reflux an azeotrope loading of F1 = 0.57 mol.min-1 and a solvent to feed ratio of Fr = 1.58 mol.mol-1 or xr = 0.42 are considered optimum over 5 Bale units.

Under countercurrent fixed bed reactor (CFBR) conditions using Bale packing, ethylene glycol promotes isobutylene solubility and improves mass transfer coefficients most likely through increased packing wetting. Mass transfer coefficients for isobutylene

Chapter 9: Conclusion and Recommendation 428 transport are most improved through higher flowrates and better Bale liquid renewal as indicated by lower Per. The average overall effectiveness factors for operation over Bale packing under conditions of recirculation are 0.11 and 0.03, for TBA and MET respectively. The selectivity towards TBA is much higher under CFBR conditions than that predicted by Equation 4.67. High concentrations of TBA inhibit the formation of MET.

Under catalytic extractive distillation (CED) conditions, TBA was enriched and the product stream was generally free of any MET formed. The entrainer, ethylene glycol, has a positive effect upon reaction rate under conditions of zero reflux and absorption type operation. Ethylene glycol improves the availablity of isobutylene and enriches the product stream. A clear improvement in TBA purity is observable, relative to catalytic distillation, in the absence of ethylene glycol. Over the range of the CED study, at -1 -1 atmospheric conditions, the rate of TBA formation, rTBA.obs (mol.g .s ), is given by:

−−09 09 −09 −09 rFTBA. obs =−1.75 × 10 + 3.1 × 10W + 1.97 × 10FEG + 9.66 × 10 yIB (7.28)

-1 The optimum reaction rate over the range of the study occurs at Fw = 0.6 mol.min , FEG -1 -09 -1 -1 = 0.61 mol.min and yIB = 0.44, corresponding to rTBA,obs = 5.84 × 10 mol.g .s .

Extractive distillation, CFBR, and CED studies were conducted in the pre-loading regime. The loading and flooding regimes for Bale packing as prepared are defined by:

−0.78 ReG, Loading = 21580ReL (5.41) and

−0.35 ReG, Flooding = 7357 ReL (5.42)

for the range 500 ≤ ReG ≤ 3000 and 10 ≤ ReL ≤ 100. Bale packing displays mixing characteristics, which are indicative of three levels of porosity reactor, with relatively poor liquid renewal at low liquid flowrates. The liquid phase backmixing in the preloading regime displays a clear minimum, which defines two regions and these are described by:

Chapter 9: Conclusion and Recommendation 429

Region 1. () 250≤≤ ReL, 623 : PerLB, = 21.59 − 0.025ReL (5.63) and

−03 Region 2. () 623 ≤≤ ReL,, 1226 : PerLR = 1.9738 +× 6.9 10 ReL (5.64)

Liquid backmixing is independent of gas flowrate in the pre-loading regime. This behaviour is considered to be a symptom of multiple levels of porosity of bale packing and the exchange between film flow along bag walls and packed flow. The liquid phase holdup of Bale packing in the preloading regime is correlated by:

−04 1.02 H0LBale =+×0.22 1 10 ReL (5.59)

The static liquid holdup of the packing was determined to be HL.S = 0.20.

Bale packing gas phase backmixing is described by:

0.19 −0.00055ReL PerGB,, =13.04ReG .10 (5.54)

Gas phase backmixing tends towards plug flow and is affected by both gas and liquid flowrates. Bale packing is recommended for operation under CFBR conditions, of relatively low gas and high liquid flowrates, where good renewal is attainable. For the purpose of catalytic distillation, it is only recommended for systems with inherently good mass transfer properties.

The requirement of higher rates of isobutylene mass transport, improved liquid renewability and greater flexibility suggests the need for improved catalytic distillation hardware. The Basket Impeller Column (BIC) is proposed as an alternative catalytic distillation reactor design. Distillative separation and reactive separation have been demonstrated to be attainable. A sieve plate design of hole size dh = 4mm and 10 %OA gives an adequate range of operation for catalytic distillation work. The rotational speed

Chapter 9: Conclusion and Recommendation 430 of the baskets gives an additional process variable. It is possible to manipulate Damköhler number using directly the speed of rotation. The BIC is recommend for use with catalytic distillation system of inherently poor mass transfer properties.

The solvent ethylene glycol, for the hydration of isobutylene to TBA under conditions of catalytic extractive distillation, warrants further investigation as it has been demonstrated that it:

• Is very effective at breaking the TBA/water azeotrope. • Enhances directly TBA reaction rate under CED conditions. • Improves TBA product purity under CED conditions relative to catalytic distillation in the absence of solvent. • Offers double the isobutylene solubility of pure water. • Improves isobutylene mass transfer coefficients. • Does react with isobutylene, however, the formation of the product MET is limited under CED conditions.

It is recommended that an evaluation of chemical equilibrium of the reaction network, formed upon the concurrent hydration and glycolisation of isobutylene over Amberlyst 15, be undertaken. This investigation should be carried out in a batch manner under pressurised conditions, given that a semi-batch mode of operation cannot be used for equilibrium determination.

In conclusion, it is recommended that the physical and thermodynamic properties of MET and DET be investigated to the extent that these chemicals can be defined in a standard simulation package. This will enable further development to be achieved using both experimental and simulation techniques. It is advised that a high proportion of the development be experimental, as the presence of side reactions has been shown to have strong non-linear interactions, which may be difficult to predict, by simulation alone. The direction of process pressurisation in the presence of a solvent should be pursued for improved isobutylene concentration, which remains the limiting factor of the process. The BIC is encouraged for further experimental development of CED for the synthesis of TBA.

Chapter 9: Conclusion and Recommendation 431 Appendix A1: Materials and Specifications

Liquid Bulk Chemicals and Standards

2-Methyl-2-Propanol, Tert-Butyl Alcohol, (TBA), [76-65-0], (Fine, used as standard), Aldrich Chemicals, 36,053-8, 99+%, reagent grade, FW 74.12, mp25-26 oC, bp83 oC, d 0.775, Warnings: Flammable liquid and Irritant.

2-Methyl-2-Propanol, Tert-Butyl Alcohol, (TBA), [76-65-0], (Bulk, used in extractive distillation experiments), APS Chemicals Limited, UNILAB, Reagent grade, A113-20L, 95 mol % minimum purity.

Ethanediol, Ethylene Glycol, (EG), [107-21-1], APS Chemicals Limited, UNILAB, Reagent Grade, 210-20L, 95 mol % minimum purity.

Ethanol, Ethyl Alcohol, (EtOH), [64-17-5], Aldrich Chemicals, 45,984-4, 99%, 1L, FW 46.07, bp78 oC, d0.790, Warnings: Flammable liquid and Irritant.

2,4,4-Trimethyl-1-Penetene, Diisobutylene, (DIB), [107-39-1], Aldrich Chemicals, T7,840-9, 99%, 5g, FW112.22, bp101-102 oC, d 0.708, Warnings: Flammable Liquid and Irritant.

Ethylene Glycol Mono-Tert-Butyl Ether, (MET), [7580-85-0], Tokyo Kasei Kogyo Co LTD, TCI_GR E0354, 25 mL, FW 118.17, bp 152 oC, Warnings: Flammable.

Sulfuric Acid, APS Chemicals, 97 to 99 wt%

Gases

2-Methylpropene, Isobutylene, (IB), [115-11-7], Matheson Gases, BOC Gases, 2.1 UN1055, mp-140, bp-6.9, Warnings: Flammable Gas. Nitrogen, Linde Gases, grade 4, Warnings: Compressed Gas. Helium, Linde Gases, grade 4.6, Warnings: Compressed Gas. Air, Linde Gases, grade instrument air, Warnings: Compressed Gas. Hydrogen, Linde Gases, grade 4, Warnings: Flammable Gas.

Catalyst

Amberlyst 15, ion exchange resin, dry, [39389-20-3], Rohm and Haas, 25 L bag, avg. size 0.5 mm, internal surface area 55 m2g-1, weight capacity 4.7 mEq H+g-1, porosity 36 %, temperature stability 120 oC. Warnings: Potentially explosive when mixed with nitric acid or strong oxidising agents. Avoid dust inhalation.

432 Appendix A2: Solubility and Kinetics Data

-1 Table A2.01 Solubility Runs, fixed: Ω = 1400 (rpm), QG = 0.2 L.min , VL = 0.35 L. +03 Run Temperature xW xEG xTBA yIB S × 10 (K) (mol.L-1) 1 315 1 0 0 1 3.358 2 335 1 0 0 0.75 2.471 3 355 1 0 0 0.5 1.800 4 315 1 0 0 1 2.396 5 335 1 0 0 0.75 2.182 6 355 1 0 0 0.5 1.644 7 315 1 0 0 1 2.185 8 335 1 0 0 0.75 1.759 9 355 1 0 0 0.5 trace 10 315 0 1 0 1 8.567 11 335 0 1 0 0.75 7.256 12 355 0 1 0 0.5 5.987 13 315 0 1 0 1 5.286 14 335 0 1 0 0.75 5.787 15 355 0 1 0 0.5 5.291 16 315 0 1 0 1 4.947 17 335 0 1 0 0.75 4.059 18 355 0 1 0 0.5 3.604 19 315 0 0 1 1 52.212 20 335 0 0 1 0.75 46.837 21 355 0 0 1 0.5 40.894 22 315 0 0 1 1 51.158 23 335 0 0 1 0.75 43.392 24 355 0 0 1 0.5 42.334 25 315 0 0 1 1 39.374 26 335 0 0 1 0.75 37.802 27 355 0 0 1 0.5 37.984 28 315 0.5 0.5 0 0.5 3.345 29 335 0.5 0.5 0 0.5 3.004 30 355 0.5 0.5 0 0.5 2.655 31 315 0.5 0 0.5 0.5 36.409 32 335 0.5 0 0.5 0.5 34.339 33 355 0.5 0 0.5 0.5 31.252 34 315 0 0.5 0.5 0.5 32.431 35 335 0 0.5 0.5 0.5 27.494 36 355 0 0.5 0.5 0.5 25.786 37 315 0.33 0.33 0.33 0.5 31.512 38 335 0.33 0.33 0.33 0.5 27.400 39 355 0.33 0.33 0.33 0.5 26.500 40 315 0.67 0.17 0.17 0.5 38.315 41 335 0.67 0.17 0.17 0.5 33.786 42 355 0.67 0.17 0.17 0.5 30.915 43 315 0.17 0.67 0.17 0.5 30.338 44 335 0.17 0.67 0.17 0.5 29.993 45 355 0.17 0.67 0.17 0.5 29.205 46 315 0.17 0.17 0.67 0.5 63.315 47 335 0.17 0.17 0.67 0.5 54.734 48 355 0.17 0.17 0.67 0.5 53.257

433 -1 Table A2.02 Internal and External Transport Evaluation (QG = 0.2 L.min , VL = 0.35 L).

08 09 Run T (K) xEG yIB dp rpm m r × 10 r × 10 (μm) (g) mol/L.s mol/g.s 49 341 0.81 0.25 450 600 3.1882 6.761 7.422 50 341 0.81 0.25 450 800 3.0319 7.362 8.499 51 341 0.81 0.25 450 1000 3.0320 7.242 8.359 52 341 0.81 0.25 450 1200 3.0162 8.570 9.944 53 341 0.81 0.25 450 1400 3.0136 8.282 9.619 54 341 0.81 0.25 450 1600 3.1012 8.500 9.593 06 07 Run T (K) xEG yIB dp rpm m r × 10 r × 10 (μm) (g) 55 361 0 0.25 350 1400 2.5267 1.233 1.708 56 361 0 0.25 350 1400 2.6982 1.405 1.822 57 361 0 0.25 450 1400 2.7670 1.172 1.482 58 361 0 0.25 450 1400 3.1013 1.335 1.506 59 361 0 0.25 550 1400 2.4898 0.790 1.111 60 361 0 0.25 550 1400 2.7151 1.058 1.363

Table A2.03 TBA Reaction Rate, Solvent Free and Solvent Mediated Hydration of -1 Isobutylene (Ω = 1400 (rpm), QG = 0.2 L.min , VL = 0.35 L, dp = 450 μm), Data presented as used in Correlation for the Development of Model M5.

Run T (K) xEG yIB CIB CW CEG rTBA (mol.L-1) (mol.L-1) (mol.L-1) (mol.g-1.s-1) 61 361 0.00 0.25 4.38E-04 55.49 0.00 3.45E-07 62 361 0.00 0.5 8.75E-04 55.49 0.00 6.93E-07 63 361 0.00 0.75 1.31E-03 55.49 0.00 9.79E-07 64 351 0.00 0.25 5.04E-04 55.49 0.00 3.46E-07 65 351 0.00 0.5 1.01E-03 55.49 0.00 4.79E-07 66 351 0.00 0.75 1.51E-03 55.49 0.00 6.33E-07 67 341 0.00 0.25 5.86E-04 55.49 0.00 1.95E-07 68 341 0.00 0.5 1.17E-03 55.49 0.00 2.63E-07 69 341 0.00 0.75 1.76E-03 55.49 0.00 4.08E-07 70 341 0.32 0.25 1.10E-03 27.02 12.45 1.06E-07 71 341 0.32 0.5 2.21E-03 27.02 12.45 1.98E-07 72 341 0.32 0.75 3.31E-03 27.02 12.45 4.15E-07 73 351 0.32 0.25 1.02E-03 27.02 12.45 1.95E-07 74 351 0.32 0.5 2.05E-03 27.02 12.45 3.65E-07 75 351 0.32 0.75 3.07E-03 27.02 12.45 5.85E-07 76 361 0.32 0.25 9.57E-04 27.02 12.45 2.30E-07 77 361 0.32 0.5 1.91E-03 27.02 12.45 4.22E-07 78 361 0.32 0.75 2.87E-03 27.02 12.45 7.58E-07 79 341 0.55 0.25 1.49E-03 12.09 14.66 3.83E-08 80 341 0.55 0.5 2.98E-03 12.09 14.66 4.52E-08 81 341 0.55 0.75 4.47E-03 12.09 14.66 1.54E-07 82 351 0.55 0.25 1.42E-03 12.09 14.66 5.34E-08 83 351 0.55 0.5 2.84E-03 12.09 14.66 1.84E-07 84 351 0.55 0.75 4.25E-03 12.09 14.66 2.80E-07 85 361 0.55 0.25 1.36E-03 12.09 14.66 7.88E-08 86 361 0.55 0.5 2.71E-03 12.09 14.66 2.21E-07 87 361 0.55 0.75 4.07E-03 12.09 14.66 5.22E-07 88 341 0.81 0.25 1.76E-03 3.95 16.41 8.43E-09 89 341 0.81 0.5 3.51E-03 3.95 16.41 1.81E-08 90 341 0.81 0.75 5.27E-03 3.95 16.41 3.00E-08 91 351 0.81 0.25 1.67E-03 3.95 16.41 9.31E-09

434 92 351 0.81 0.5 3.34E-03 3.95 16.41 2.29E-08 93 351 0.81 0.75 5.01E-03 3.95 16.41 1.22E-07 94 361 0.81 0.25 1.60E-03 3.95 16.41 3.60E-08 95 361 0.81 0.5 3.19E-03 3.95 16.41 9.83E-08 96 361 0.81 0.75 4.79E-03 3.95 16.41 2.32E-07 97 361 0.00 0.75 1.31E-03 55.49 0.00 9.89E-07 98 361 0.11 0.75 1.78E-03 43.74 5.18 9.02E-07 99 361 0.20 0.75 2.24E-03 32.78 8.10 9.23E-07 100 361 0.32 0.75 2.87E-03 27.02 12.45 7.58E-07 101 361 0.37 0.75 3.17E-03 20.29 11.95 8.73E-07 102 361 0.47 0.75 3.71E-03 14.60 13.01 7.64E-07 103 361 0.55 0.75 4.07E-03 12.09 14.66 5.22E-07 104 361 0.64 0.75 4.44E-03 8.08 14.53 5.53E-07 105 361 0.73 0.75 4.68E-03 5.62 15.43 3.39E-07 106 361 1.00 0.75 4.624E-03 0.00 17.95 0.00E+00

Table A2.05 MET Reaction Rate, Water Mediated Glycolisation of Isobutylene (Ω = 1400 -1 (rpm), QG = 0.2 L.min , VL = 0.35 L, dp = 450 μm). Run T (K) xEG yIB CIB CW CEG rMET (mol.L-1) (mol.L-1) (mol.L-1) (mol.g-1.s-1) 88 341 0.81 0.25 1.76E-03 3.95 16.41 9.54E-07 89 341 0.81 0.50 3.51E-03 3.95 16.41 2.34E-06 90 341 0.81 0.75 5.27E-03 3.95 16.41 4.10E-04 91 351 0.81 0.25 1.67E-03 3.95 16.41 1.38E-06 92 351 0.81 0.50 3.34E-03 3.95 16.41 3.43E-06 93 351 0.81 0.75 5.01E-03 3.95 16.41 4.68E-06 94 361 0.81 0.25 1.60E-03 3.95 16.41 1.99E-06 95 361 0.81 0.50 3.19E-03 3.95 16.41 4.22E-06 96 361 0.81 0.75 4.79E-03 3.95 16.41 8.72E-06 97 361 0.00 0.75 1.31E-03 55.49 0.00 0.00E+00 98 361 0.11 0.75 1.78E-03 43.74 5.18 0.00E+00 99 361 0.20 0.75 2.24E-03 32.78 8.10 0.00E+00 100 361 0.32 0.75 2.87E-03 27.02 12.45 0.00E+00 101 361 0.37 0.75 3.17E-03 20.29 11.95 9.35E-07 102 361 0.47 0.75 3.71E-03 14.60 13.01 1.38E-06 103 361 0.55 0.75 4.07E-03 12.09 14.66 0.00E+00 104 361 0.64 0.75 4.44E-03 8.08 14.53 2.55E-06 105 361 0.73 0.75 4.68E-03 5.62 15.43 3.57E-06 106 361 1.00 0.75 4.62E-03 0.00 17.95 8.58E-06

435 Appendix A3: Digital Image Processing

The software IrfanView was used to batch process the images obtained converting them to the tiff format. IrfanView, Author: Irfan Skiljan, http://www.irfanview.com

These images were then processed in ScionImage, which allowed for measurement of green light intensity. The following is a program developed to batch process photos digital photos within ScionImage. The language used is specific to the ScionImage compiler and is based upon Pascal. The complier did not accept loop programming at the time. The following example processes photos 1 to 2 while excluding photo 3 which is distinguished by use of { } brackets. The pixel coordinates used to define the sample window need to be obtained before running the program (x and y refer to the left top corner of the sample window, w and h are its width and height). Slice N = 2 refers to the green light split of the image.

macro 'batchm'; VAR IM:string; i,N,M:integer; x,y,w,h:real; BEGIN x:=GetNumber('Enter x:',0); y:=GetNumber('Enter y:',0); w:=GetNumber('Enter w:',0); h:=GetNumber('Enter h:',0); N:=2; OPEN('c:\calc\IM001.tif'); SelectSlice(N); MakeRoi(x,y,w,h); Measure; Save; Close; OPEN('c:\calc\IM002.tif'); SelectSlice(N); MakeRoi(x,y,w,h); Measure; Save; Close; {OPEN('c:\calc\IM003.tif'); SelectSlice(N); MakeRoi(x,y,w,h); Measure; Save; Close;} End;

436 Appendix A4: RTD Data Smoothing and Processing

The raw RTD data obtained was processed in Microsoft Excel using the following macros, which were specifically developed for the purpose. The macros are written in Visual Basic for Applications, using the Visual Basic Editor available in Excel. The spreadsheet in which the macros have been written is kept open in the background allowing for general access of the macros. The macros are called from the Tools menu or by predefined function keys. The subject data is first highlighted in Excel. The first cell of the highlighted column of data forms the (1,1) entry of the matrix defining the space (i,j) of the macro. The macros included are those for polynomial smoothing of window size 5 and 9, and determinations of estimates of mean residence time and standard deviation using the method of moments.

VBA Macro code:

Option Explicit

‘Savitzky-Golay Polynomial Smoothing Window Size 5 Sub smoothing05() Dim i, j, N, Nrow, p As Integer, c, m As Double Nrow = Selection.Rows.Count N = InputBox("What is n?") p = 3 For j = 1 To N For i = p To (Nrow - p + 1) c = -3 * (Selection.Cells(i + 2, j)) + 12 * (Selection.Cells(i + 1, j)) + 17 * (Selection.Cells(i, j)) + 12 * (Selection.Cells(i - 1, j)) - 3 * (Selection.Cells(i - 2, j)) m = c / 35 Selection.Cells(i, j + 1) = m Next i p = p + 2 Next j End Sub

‘Savitzky-Golay Polynomial Smoothing Window Size 9 Sub aasmoothing09() Dim i, j, N, Nrow, p As Integer, c, m As Double Nrow = Selection.Rows.Count N = InputBox("What is n?") p = 5 For j = 1 To N For i = p To (Nrow – p + 1) c = -21 * (Selection.Cells(i + 4, j)) + 14 * (Selection.Cells(i + 3, j)) + 39 * (Selection.Cells(i + 2, j)) + 54 * (Selection.Cells(i + 1, j)) + 59 * (Selection.Cells(i, j)) + 54 * (Selection.Cells(i - 1, j)) + 39 * (Selection.Cells(i - 2, j)) + 14 * (Selection.Cells(i - 3, j)) - 21 * (Selection.Cells(i - 4, j)) m = c / 231 Selection.Cells(i, j + 1) = m Next i

437 p = p + 4 Next j End Sub

‘Simpson Integration Sub Simpson(j As Integer, N As Integer, thn As Integer) Dim sum As Double, i, l As Integer, a As Double sum = 0 For i = thn + 2 To (N - 1) If i Mod 2 = 0 Then sum = sum + 4 * Selection.Cells(i, j) Else: sum = sum + 2 * Selection.Cells(i, j) End If Next i a = (sum + Selection.Cells(thn + 1, j) + Selection.Cells(N, j)) * ((1 / 30) / 3) Selection.Cells(N + 2, j) = a End Sub

‘Main Data Analysis Macro, Method of Moments Sub aamain() Dim v As Double, B As Double, Bavg As Double, Diff As Double, tm As Double, ai As Double, sum As Double, sumI As Double, Iavg As Double, l As Integer, j As Integer, i As Integer, th As Integer, thn As Integer, Nrow As Integer, g1 As Double, g2 As Integer, g3 As Double, g4 As Integer, Nnew As Integer, scan As Integer, k As Integer, a As Double, X As Double, Y As Double, E As Double

‘threshold range input and adjustment of baseline Nrow = Selection.Rows.Count th = InputBox("What is th?") thn = InputBox("what is thn?") scan = InputBox("what is scan?") B = 0 For i = th To thn B = B + Selection.Cells(i, 2) Next i Bavg = B / (thn - th + 1) Selection.Cells(Nrow + 3, 2) = Bavg For i = 1 To thn Selection.Cells(i, 3) = 0 Next i For i = thn + 1 To Nrow Diff = Selection.Cells(i, 2) - Bavg If Diff < 0 Then Diff = 0 ElseIf Diff >= 0 Then Diff = Diff End If Selection.Cells(i, 3) = Diff Next i Nnew = scan - 1 For i = scan To Nrow If (Selection.Cells(i, 3) = 0) Then Exit For

438 Else Nnew = Nnew + 1 End If Next i If (Nnew Mod 2 = 0) Then Nnew = Nnew - 1 End If

‘calculation of area under the curve Simpson 3, Nnew, thn 'E a = Selection.Cells(Nnew + 2, 3) Selection.Cells(1, 9) = Selection.Cells(Nnew + 2, 3) For i = 1 To Nnew Selection.Cells(i, 4) = Selection.Cells(i, 3) / a Next i

'tm For i = 1 To Nnew E = Selection.Cells(i, 4) Selection.Cells(i, 5) = E * Selection.Cells(i, 1) Next i Simpson 5, Nnew, thn tm = Selection.Cells(Nnew + 2, 5) Selection.Cells(2, 9) = Selection.Cells(Nnew + 2, 5)

't -tm For i = 1 To Nnew Selection.Cells(i, 6) = Selection.Cells(i, 1) - tm Next i

'standard deviation For i = 1 To Nnew Selection.Cells(i, 7) = (Selection.Cells(i, 6) ^ 2) * Selection.Cells(i, 4) Next i Simpson 7, Nnew, thn Selection.Cells(3, 9) = Selection.Cells(Nnew + 2, 7)

'skewness For i = 1 To Nnew Selection.Cells(i, 8) = ((Selection.Cells(i, 6) ^ 3) * Selection.Cells(i, 4)) * (1 / (Selection.Cells(Nnew + 2, 7) ^ 3 / 2)) Next i Simpson 8, Nnew, thn Selection.Cells(4, 9) = Selection.Cells(Nnew + 2, 8)

End Sub

439 Appendix A5: UNIFAC Based VLE Calculation

The following Tables demonstrate the VLE calculation used, based upon UNIFAC, for an example composition of x’TBA = 0.65 with xr = 0.5. The tables are taken directly from a spreadsheet used to perform the calculation. Values entered such as composition, interaction parameters, and constants are depicted in bold, while final results are shown in italics.

VLE Calculation: Bubble Point Calculation

Component id and molar fraction ID Component x 1 TBA 0.325 2 WATER 0.175 3 EG 0.5 4 IB 0

Antoine Constants as employed in Equation (6.18) Comp. A1 A2 A3 TBA 1 16.8548 2658.2900 -95.5000 H2O 2 18.3036 3816.4400 -46.1300 EG 3 20.2501 6022.1800 -28.2500 IB 4 16.9483 2857.7498 11.9411

Temperature estimated at P = 760 mmHg based upon molar contribution, equation (6.18) T1 355.5690 T2 373.1521 T3 470.5116 T4 265.1074

Tinitial 416.1173

Pressure estimation based upon Tinitial Equation (6.18) P1 5237.7065 P2 2947.0042 P3 112.5851 P4 28918.1897

Activity Coefficients as calculated by UNIFAC in next Section ac1 1.4745 ac2 1.6327 ac3 1.0546 ac4 5.4225

New pressure estimate Equation (6.17) P1 new 1166.8675 P2 new 656.5399 P3 new 25.0819 P4 new 6442.4565

New Temperature estimate Equation (6.18) and Temperature Iteration Loop T1 new 366.9557 T2 new 369.1023 T3 new 381.9144

440 T4 new 337.5181

T new 374.8107 T in = T new 374.8107 T final 374.8107

Result: Vapour Composition yi based upon modified Rault’s Law Equation (6.16) y1 0.7358 y2 0.2468 y3 0.0174 y4 0.0000

Result: Vapour and Liquid Compositions Solvent Free Basis 1 2 TBA W x 0.65 0.35 y 0.75 0.25

Result: Relative Volatility at x’TBA = 0.65 and xr = 0.5 α 1.605

Activity Coefficients Calculation using UNIFAC, Smith et al (1996) Chapter 6:

UNIFAC VLE subgroup parameters (AspenPlus) C CH CH2 CH3 OH H2O 1000 1005 1010 1015 1200 1300 Q 0 0.228 0.54 0.848 1.2 1.4 R 0.2195 0.4469 0.6744 0.9011 1 0.92

Component Group Breakdown 1000 1005 1010 1015 1200 1300 vk1 1 0031 0 vk2 0 0000 1 vk3 0 0202 0 vk4 1 0120 0

Calculation of ri and qi by Equation (6.09) and (6.10) ii 1 2 3 4 q 3.744 1.4 3.48 2.236 r 3.9228 0.92 3.3488 2.6961

UNIFAC VLE Interaction parameters amk (K) (AspenPlus) C CH CH2 CH3 OH H2O amk 1000 1005 1010 1015 1200 1300 1000 0 0 0 0 986.5 1318 1005 0 0 0 0 986.5 1318 1010 0 0 0 0 986.5 1318 1015 0 0 0 0 986.5 1318 1200 156.4 156.4 156.4 156.4 0 353.5 1300 300 300 300 300 -229.1 0

Temperature Input from VLE calculation above. Temp VLE 374.8107

441

Calculation of τmk by Equation (6.15)

τmk 1000 1005 1010 1015 1200 1300 1000 1.0000 1.0000 1.0000 1.0000 0.0719 0.0297 1005 1.0000 1.0000 1.0000 1.0000 0.0719 0.0297 1010 1.0000 1.0000 1.0000 1.0000 0.0719 0.0297 1015 1.0000 1.0000 1.0000 1.0000 0.0719 0.0297 1200 0.6588 0.6588 0.6588 0.6588 1.0000 0.3894 1300 0.4491 0.4491 0.4491 0.4491 1.8427 1.0000

Calculation of eik by Equation (6.11) eik 1000 1005 1010 1015 1200 1300 1 0.0000 0.0000 0.0000 0.6795 0.3205 0.0000 2 0.0000 0.0000 0.0000 0.0000 0.0000 1.0000 3 0.0000 0.0000 0.3103 0.0000 0.6897 0.0000 4 0.0000 0.0000 0.2415 0.7585 0.0000 0.0000

Calculation of βik by Equation (6.12)

βik 1000 1005 1010 1015 1200 1300 1 0.8907 0.8907 0.8907 0.8907 0.3694 0.1450 2 0.4491 0.4491 0.4491 0.4491 1.8427 1.0000 3 0.7647 0.7647 0.7647 0.7647 0.7120 0.2778 4 1.0000 1.0000 1.0000 1.0000 0.0719 0.0297

Calculation of θk and sk by Equation (6.13) and (6.14) 1000 1005 1010 1015 1200 1300

θk 0.0000 0.0000 0.1687 0.2582 0.4966 0.0765 sk 0.7884 0.7884 0.7884 0.7884 0.6683 0.2826

Calculation Ji and Li by Equation (6.07) and (6.08) 1 2 3 4

Ji 1.2612 0.2958 1.0767 0.8668 Li 1.1693 0.4373 1.0869 0.6984 ln yci 0.0258 -0.0425 -0.0020 0.2712

Internal Calculation of Summation within Equation (6.08) (to obtain Li) comp 1000 10051010 1015 1200 1300 1 0 0 0.1905 0.2089 0.4645 0.0393 2 0 0 0.0961 0.1471 1.3693 -0.9930 3 0 0 0.1731 0.2505 0.4854 0.0752 4 0 0 0.1565 0.1472 0.0535 0.0080

Calculation of activity coefficient γi Equation (6.06), and Equation (6.04) COMP: 1 2 3 4 ln yri 0.3625 0.5328 0.0552 1.4194 ln γci 0.3883 0.4903 0.0532 1.6906

γi 1.4745 1.6327 1.0546 5.4225

442 Appendix A6: Wilson Model Based VLE Calculation

The following Tables demonstrate the VLE calculation used, based upon Wilson model, for an example composition of x’TBA = 0.65 with xr = 0.5. The tables are taken directly from a spreadsheet used to perform the calculation. Values entered such as composition, interaction parameters, and constants are depicted in bold, while final results are shown in italics.

VLE Calculation: Bubble Point Calculation

Component id and molar fraction ID Component x 1 TBA 0.325 2 WATER 0.175 3 EG 0.5 4 IB 0

Antoine Constants as employed in Equation (6.18) Comp. A1 A2 A3 TBA 1 16.8548 2658.2900 -95.5000 H2O 2 18.3036 3816.4400 -46.1300 EG 3 20.2501 6022.1800 -28.2500

Temperature estimated at P = 760 mmHg based upon molar contribution, equation (6.18) T1 355.5690 T2 373.1521 T3 470.5116

Tinitial 416.1173

Pressure estimation based upon Tinitial Equation (6.18) P1 5237.7065 P2 2947.0042 P3 112.5851

Activity Coefficients as calculated by UNIFAC in next Section ac1 1.6980 ac2 0.6338 ac3 0.8565

New pressure estimate Equation (6.17) P1 new 1219.0221 P2 new 685.8848 P3 new 26.2030

New Temperature estimate Equation (6.18) and Temperature Iteration Loop T1 new 368.1732 T2 new 370.3019 T3 new 382.8249

T new 375.8716 T in = T new 375.8716 T final 375.8716

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Result: Vapour Composition yi based upon modified Rault’s Law Equation (6.16) y1 0.8851 y2 0.1001 y3 0.0148

Result: Vapour and Liquid Compositions Solvent Free Basis 1 2 TBA W x 0.65 0.35 y 0.90 0.10

Result: Relative Volatility at x’TBA = 0.65 and xr = 0.5 α 1.605

Activity Coefficients Calculation using the Wilson model:

Wilson Binary Interaction Coefficients, λij, Liu et al (1993) Chapter 6

λij 1 2 3 1 7402.617 -379.009 2 5419.225 4143.67 3 3627.66 -7460.31

Temperature Input from VLE calculation above. T 375.8716

Molar Volume Coefficients of Equation (6.03) a1 0.000132 1.01E-05 3.55E-05 a0 0.055463 0.014994 0.045172

Calculation of Molar Volumes by Equation (6.03) 1 2 3 V 0.105228 0.018799 0.058518

Calculation of Wilson Parameters by Equation (6.02) Λij 1 2 3 1 1 0.01672 0.627812 2 0.988228 1 0.826578 3 0.563237 3.496583 1

Calculation of Activity coefficient γi by Equation (6.01) 1 2 3 ln γ 0.529436 -0.45607 -0.15491

γi 1.697974 0.633772 0.856491

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