INFORMATION TO USERS

This manuscript has been reproduced from the microfilm master. UMI films the text directly from the original or copy submitted Thus, some thesis and dissertation copies are in typewriter face, while others may be from any type of computer printer.

The quality of this reproduction is dependent upon the quality of the copy submitted. Broken or indistinct print, colored or poor quality illustrations and photographs, prim bleedthrough, substandard margins, and improper alignment can adversely affect reproduction.

In the unlikely event that the author did not send UMI a complete manuscript and there are missing pages, these will be noted. Also, if unauthorized copyright material had to be removed, a note wiD indicate the deletion.

Oversize materials (e.g^ maps, drawings, charts) are reproduced by sectioning the original, beginning at the upper left-hand comer and continuing from left to right in equal sections with small overlaps. Each original is also photographed in one exposure and is included in reduced form at the back of the book.

Photographs included in the original manuscript have been reproduced xerographically in this copy. Higher quality 6" x 9" black and white photographic prints are available for any photographs or illustrations appearing in this copy for an additional charge. Contact UMI directly to order.

A Beu & Howell information Company 300 Norm ZeeD ftoao Ann Arbor. Ml 48105-1346 USA 313/761-4700 600/521-0600 A NOVEL IMMOBILIZED CELL BIOREACTOR AND MEMBRANE

ULTRAFILTRATION FOR XANTHAN GUM PRODUCTION

DISSERTATION

Presented in Partial Fulfillment of the Requirements for

the Degree Doctor of Philosophy in the Graduate

School of The Ohio State University

By

Yang-Ming Lo, B.S., M.S.

The Ohio State University

1995

Dissertation Committee: Approved by

D.B. Min

S.T. Yang Co-Advisor

J.H. Litchfield

Q.H. Zhang Co-Advisor

Food Science and Nutrition UMI Number: 9544624

UMI Microform 9544624 Copyright 1995, by UMI Company. All rights reserved.

This microform edition is protected against unauthorized copying under Title 17, United States Code.

UMI 300 Nbrth Zeeb Road Ann Arbor, MI 48103 To my Family

ii ACKNOWLEDGEMENTS

I express sincere appreciation to my co-advisors, Dr. Shang-Tian Yang, for his invaluable guidance, inspiration, insight, and encouragement throughout my research and coursework in Biochemical Engineering, and Dr. David B. Min, for his generous advice and support that lead me into fermentation technology, a combination of Food Science and Chemical Engineering. I would like to express my gratitude to the Ohio Com

Growers Association and Midwest Advanced Food Manufacturing Alliance for their financial support. Thanks go to the other members of my advisory committee, Drs. John

H. Litchfield and Qinghua H. Zhang, for their suggestions and comments. Gratitude is expressed to the groups of biochemical engineering and polymers in the Department of

Chemical Engineering for their support and facilities.

The technical assistance of Michael Kukla, Carl Scott, Don Ordaz, Kathy Wolken, and John Mitchell is gratefully acknowledged. I would like to thank Jyh-Wen Yen,

Chen-Shcn Wu, Ellen Silva, Hui Zhu, Jin-Jae Chen, Chin-Yao Huang, Wen-Lin Tsai, and

Donald Streibig for their cordial friendship and encouragement. 1 am deeply grateful to

Su-Tai Wang for giving me an immediate lift to campus on the occasion that my car was dead right before my oral defense without any warning. To my mother and two younger brothers, Yang-Ping and Yang-Tai, I thank you for your understanding of my absence for the past four years and for your enormous support when times were rough for our family.

To my wife, Nien-Chen, I offer sincere thanks for your unshakable faith in me and your willingness to endure with me the vicissitudes of my endeavors. VITA

January 13, 1967 ...... Bom - Taipei, Taiwan

June 1989 ...... B.S./Animal Science National Taiwan University Taipei, Taiwan

1989-1991 ...... Second Lieutenant, Infantry Taitung, Taiwan

August 1993 ...... M.S./Food Science and Nutrition The Ohio State University Columbus, Ohio

1993-Present ...... Graduate Research Associate The Ohio State University Columbus, Ohio

FIELDS OF STUDY

Major Field: Food Science and Technology

Studies in Food Chemistry, Food Engineering, and Biochemical Engineering TABLE OF CONTENTS

ACKNOWLEDGEMENT...... iii

VITA ...... iv

LIST OF TABLES...... xni

LIST OF FIGURES...... xiv

ABSTRACT...... xviii

CHAPTER PAGE

I. INTRODUCTION...... 1

1.1 Background ...... 1

1.1.1 Xanthan gum fermentation ...... 1

1.1.2 Immobilized cell bioneactor ...... 3

1.1.3 Ultrafiltration of xanthan solution ...... 6

1.2 Objectives ...... 7

1.3 Scope of the Study ...... 9

II. LITERATURE REVIEW ...... 13

2.1 Introduction ...... 13

2.2 Structure and Conformation of Xanthan ...... 14

2.3 Characteristic Properties of Xanthan Gum ...... 18

2.3.1 Pseudoplasticity ...... 18

2.3.2 Molecular weight ...... 20

2.4 Microbial Biosynthesis of Xanthan Gum ...... 22

2.4.1 Function and importance ...... 22 2.4.2 Enzyme system ...... 23

2.4.3 Mechanism of biosynthesis ...... 24

2.5 Fermentation Kinetics ...... 26

2.5.1 Cell growth kinetics ...... 26

2.5.2 Product formation kinetics ...... 29

2.5.3 Effects of environmental factors ...... 30

2.5.3.1 Medium composition ...... 31

2.5.3.2 Temperature ...... 31

2.5.3.3 pH ...... 31

2.5.3.4 Shear ...... 32

2.5.3.5 Dissolved oxygen ...... 32

2.5.3.6 Other factors ...... 33

2.6 Oxygen Transfer in Xanthan Fermentation...... 33

2.6.1 Determination of volumetric mass transfer coefficient ...... 33

2.6.2 Factors affecting oxygen transfer ...... 35

2.7 Xanthan Production Process ...... 39

2.7.1 Bioreactor types ...... 41

2.7.2 Fermentation processes ...... 45

2.7.3 Recovery and purification ...... 46

2.7.3.1 Alcohol precipitation ...... 47

2.7.3.2 Freezing-thawing method ...... 48

2.7.3.3 Other purification methods ...... 49

2.8 Cell Immobilization ...... 50

vi 2.8.1 Advantages and limitations ...... 50

2.8.1.1 Advantages ...... 50

2.8.1.2 Limitations ...... 51

2.8.1.3 Changes in cell properties ...... 52

2.8.2 Methods of immobilization...... 53

2.8.2.1 Adsorption ...... 54

2.8.2.2 Entrapment ...... 56

2.8.2.3 Coupling ...... 56

2.8.3 Immobilized cell bioreactors ...... 56

2.9 Ultrafiltration Processes ...... 59

2.9.1 Significance and applications ...... 59

2.9.2 Theory ...... 63

2.9.3 Factors affecting ultrafiltration ...... 67

2.9.3.1 Physico-chemical factors ...... 67

2.9.3.2 Feed concentration ...... 68

2.9.3.3 Shear rate ...... 68

2.9.3.4 Viscosity ...... 68

2.9.3.5 Transmembrane pressure ...... 69

2.9.3.6 Temperature ...... 69

2.9.3.7 Other factors ...... 69

2.9.4 Membrane selection ...... 70 in. IMMOBILIZED CELL XANTHAN GUM FERMENTATION ...... 74

3.1 Summary...... 74

vii 3.2 Introduction ...... 75

3.3 Materials and Methods ...... 78

3.3.1 Culture and m edia ...... 78

3.3.2 Centrifugal, packed-bed bioreactor (CPBR) ...... 79

3.3.2.1 Bioreactor construction ...... 79

3.3.2.2 Selection of the support matrix for immobilization 82

3.3.2.3 Bioreactor startup ...... 82

3.3.2.4 Repeated batch fermentation ...... 83

3.3.2.5 Determination of immobilized cell density ...... 84

3.3.2.6 Determination of cell viability ...... 84

3.3.2.7 Scanning electron microscopy ...... 85

3.3.3 Analytical methods ...... 85

3.3.3.1 Cell density ...... 85

3.3.3.2 Glucose concentration ...... 86

3.3.3.3 Xanthan gum concentration ...... 86

3.4 Results and Discussion ...... 87

3.4.1 Fibrous matrix selection ...... 87

3.4.2 Reactor startup ...... 87

3.4.3 Liquid-continuous fermentation at 150 rpm ...... 91

3.4.4 Liquid-continuous fermentation at 350 rpm ...... 91

3.4.5 Gas-continuous fermentation ...... 95

3.4.6 Cell immobilization ...... 97

3.4.7 Free cell batch fermentation in STR ...... 100

viii 3.4.8 Comparison between STR and CPBR ...... 105

3.5 Conclusion and Recommendation ...... 105

3.6 References ...... 108

IV. OXYGEN TRANSFER IN VISCOUS XANTHAN BROTHS...... 112

4.1 Summary...... 112

4.2 Introduction...... 113

4.3 Materials and Methods ...... 118

4.3.1 Xanthan fermentation broth ...... 118

4.3.2 Stirred tank reactor (STR) ...... 118

4.3.3 C PB R ...... 120

4.3.4 Oxygen transfer experiment ...... 120

4.3.5 Determination of C* and C l ...... 121

4.3.6 Determination of kLa ...... 121

4.4 Results and Discussion ...... 122

4.4.1 C* in xanthan broth ...... 122

4.4.2 kLainSTR-DT ...... 122

4.4.3 kLa in STR-MP ...... 126

4.4.4 kt a in STR-WIO ...... 126

4.4.5 kLa in CPBR-LC ...... 127

4.4.6 ki,a in CPBR-GC ...... 128

4.4.7 Comparison of various bioreactor systems ...... 129

4.5 Conclusion and Recommendation ...... 132

4.6 References ...... 133

ix V. ULTRAFILTRATION OF XANTHAN FERMENTATION BROTH...... 136

5.1 Summary...... 136

5.2 Introduction...... 137

5.3 Experiments ...... 140

5.3.1 Xanthan broth ...... 140

5.3.2 Ultrafiltration system ...... 140

5.3.3 Process feasibility study ...... 142

5.3.4 Kinetic study ...... 143

5.3.4.1 Membrane fouling ...... 143

5.3.4.2 Effects of transmembrane pressure ...... 143

5.3.4.3 Effects of xanthan concentration ...... 143

5.3.4.4 Effects of pumping or shear rate ...... t44

5.3.4.5 Effects of pH ...... 144

5.3.5 Effects of shear on xanthan polymer ...... 144

5.3.6 Analytical methods ...... 145

5.3.6.1 Determination of xanthan concentration...... 145

5.3.6.2 Measurement of viscosity ...... 145

5.3.7 Scanning electron microscopy ...... 146

5.4 Results and Discussion ...... 146

5.4.1 Feasibility study ...... 146

5.4.1.1 Recovery yield ...... 146

5.4.1.2 Concentration polarization ...... 148

5.4.1.3 Membrane fouling ...... 150

x 5.4.2 Kinetic studies ...... 150

5.4.2.1 Effect of pressure ...... 153

5.4.2.2 Effect of xanthan concentration ...... 155

5.4.2.3 Effect of pumping (shear) rate ...... 158

5.4.2.4 Effect of solution pH ...... 160

5.4.2.5 Filtrate flux ...... 162

5.4.3 Effects of ultraflltration on xanthan molecule ...... 164

5.4.3.1 Rheological properties ...... 164

5.4.3.2 Molecular weight ...... 164

5.4.4 Energy consumption ...... 166

5.5 Conclusion ...... 168

5.6 List of Symbols ...... 169

5.7 References ...... 170

VI. ULTRAFILTRATION PROCESS EVALUATION...... 172

6.1 Summary...... 172

6.2 Introduction ...... 173

6.3 Material and Methods ...... 174

6.3.1 Xanthan broth ...... 174

6.3.2 Ultraflltration ...... 174

6.3.3 Pressure drop ...... 175

6.3.4 Rheology measurement ...... 178

6.4 Results and Discussion ...... 178

6.4.1 Ultraflltration ...... 178

xi 6.4.2 Flow behavior of xanthan solution ...... 179

6.4.2.1 Rheological characteristics ...... 179

6.4.2.2 Pressure drop for flow in tube ...... 181

6.4.3 Energy consumption for ultrafiltration ...... 187

6.4.3.1 Power consumption for pumping ...... 187

6.4.3.2 Energy consumption ...... 187

6.4.3.3 Effect of shear (pumping) rate ...... 187

6.4.3.4 Effect of transmembrane pressure ...... 189

6.4.4 Process economics ...... 189

6.4.4.1 Operating costs for ultraflltration ...... 189

6.4.4.2 Energy costs for alcohol distillation ...... 192

6.4.4.3 Total xanthan recovery costs ...... 192

6.5 Conclusions ...... 195

6.6 Notation ...... 195

6.7 References ...... 197

VII. CONCLUSIONS AND RECOMMENDATIONS...... 2(X)

7.1 CPBR for Xanthan Gum Production ...... 200

7.2 Ultraflltration of Xanthan Gum Fermentation Broth ...... 201

BIBLIOGRAPHY...... 203

APPENDICES...... 218

A. Fermentation Analytical Method ...... 218 B. Equation Derivation ...... 226 C. Experimental Data for Chapter III ...... 234 D. Experimental Data for Chapter IV ...... 243 E. Experimental Data for Chapter V ...... 247 F. Experimental Data for Chapter VI ...... 254 xii UST OF TABLES

TABLE PAGE

2.1 Solubility of oxygen at 1 atm pure oxygen in water at various temperatures, and solutions of salt or acid at 25°C ...... 36

2.2 Comparison of different bioreactor systems for xanthan gum fermentation ...... 42

2.3 Characteristics of some membrane processes ...... 60

2.4 Typical ultraflltration membrane materials ...... 71

2.5 Membrane module concepts and their characteristics...... 73 UST OF FIGURES

FIGURE PAGE

1.1 Conventional process for xanthan gum production ...... 2

1.2 Proposed process for xanthan gum production ...... 8

1.3 Characteristics of the centrifugal, packed bed bioreactor (CPBR) ...... 11

2.1 Repeated unit of xanthan gum molecules ...... 16

2.2 Conformational forms of xanthan in solution ...... 17

2.3 Proposed pathway of xanthan synthesis in Xanthomonas campestris ...... 25

2.4 A typical batch fermentation for xanthan gum production by Xanthomonas campestris (30°C)...... 27

2.5 A plot of the ln(C*-Cj) against time of aeration, the slope of which equals -*i a ...... 37

2.6 Summary of cell immobilization techniques ...... 55

2.7 Examples of components separated by MF, UF, and RO processes ...... 62

2.8 Schematic diagram and concentration profile of UF process ...... 64

2.9 Structure of polysulfone ...... 71

3.1 Schematic diagram of the centrifugal, packed-bed bioreactor: (a) reactor system; (b) configuration of the fibrous packing; and (c) fluid flow inside the bioreactor ...... 80

3.2 Cell adsorption to various fibrous matrices during xanthan fermentation: (a) 100% cotton (towel with looping); (b) 100% cotton (plane sheet without looping); (c) 50% cotton + 50% polyester; and (d) 100% polyester ...... 88

3.3 Scanning electron micrograph of Xanthomonas campestris cells adsorbed on various fibrous matrices ...... 89

3.4 Typical kinetics of two-stage, repeated-batch xanthan gum fermentation in the CPBR ...... 90

xiv 3.5 Kinetics of repeated-batch xanthan fermentation in CPBR at 150 rpm ...... 92

3.6 Xanthan yield (Yp/s) and production rate (dP/dt) from the repeated- batch fermentation in CPBR at 150 rpm ...... 93

3.7 Kinetics of repeated-batch xanthan fermentation in liquid-continuous CPBR (CPBR-LC)...... 94

3.8 Xanthan yield (Yp/s), production rate (dP/dt), and quality (in terms of the apparent viscosity of 1.8% xanthan solution) from the repeated-batch fermentation in CPBR-LC ...... 96

3.9 Kinetics of repeated-batch xanthan fermentation in gas-continuous CPBR (CPBR-GC); liquid volume was reduced from 5.0 L to 2.5 L after first batch ...... 98

3.10 Xanthan yield (Yp/s), production rate (dP/dt), and quality (in terms of their apparent viscosity of 1.8% xanthan solution) from the repeated- batch fermentation in CPBR-GC ...... 99

3.11 Scanning electron micrographs: (a) the configuration of towel matrix; (b) the attachment of Xanthomonas campestris cells on the fiber surfaces; (c) matrix sample from outer packing layer; and (d) matrix sample from inner packing layer...... 101

3.12 Kinetics of batch xanthan fermentations in stirred tank reactor (STR) with 2.5% glucose ...... 102

3.13 Kinetics of batch xanthan fermentations in stirred tank reactor (STR) with 5.0% glucose ...... 103

3.14 Comparison of (a) volumetric xanthan productivity; and (b) specific xanthan productivity of batch fermentations in STR, CPBR-LC, and CPBR-GC...... 106

3.15 Specific xanthan productivity as a function of specific oxygen uptake rate during xanthan gum fermentation ...... 107

4.1 Schematic diagrams and flow patterns of bioreactors studied: (a) STR-DT; (b) STR-MP; (c) STR-WIO; and (d) CPBR ...... 119

4.2 Typical dissolved oxygen tension profiles obtained with water and xanthan broth in the static gassing-out experiment ...... 123

4.3 Effect of xanthan concentration on oxygen solubility (C+) in xanthan broth at 30 °C...... 124 xv 4.4 kLa in various systems: (a) STR-DT; (b) STR-MP; (c) STR-WIO; and (d) CPBR-LC...... 125

4.5 in CPBR-GC at various xanthan concentrations: (a) 0%; (b) 1.2%; (c) 2.2%; and

4.6 Comparison of maximal kLa and OTR in various bioreactor systems ...... 131

5.1 Hollow-Fiber Ultraflltration system used in this study ...... 141

5.2 Unsteady-state performance of ultraflltration of xanthan broth at 25 psig and a fixed pumping rate of 3.5 (-350 mL/min) ...... 147

5.3 Steady-state performance of ultrafiltration of xanthan broths at 25 psig and -350 mLVmin pumping rate ...... 149

5.4 Effect of cells on membrane fouling ...... 151

5.5 SEM photographs of membrane before and after use in ultraflltration of xanthan broth ...... 152

5.6 Effects of pressure on filtrate flux and resistance, (a) water (b) 2.5% xanthan solution ...... 154

5.7 Effect of xanthan concentration on solution viscosity measured at (a) 500 s 1 and (b) 6 s 1 shear rate ...... 156

5.8 Effects of xanthan concentration on (a) total volumetric flow rate through the hollow-fiber ultrafilter; (b) filtrate rate; and (c) filtrate flux and fouling resistance. The transmembrane pressure was 27 psig ...... 157

5.9 Effect of shear rate on the viscosity of xanthan solution ...... 159

5.10 Effects of the pumping rate (pump speed) on (a) total volumetric flow rate through the ultrafilter; (b) filtrate rate; and (c) filtrate flux and fouling resistance, (with 3.5% xanthan solution at 27 psig) ...... 161

5.11 Effect of the pH on (a) solution viscosity; and (b) filtrate rate (at 3.5 pumping rate and 25 psig) ...... 163

5 .12 Comparison of Theological characteristics of xanthan solutions before and after ultraflltration ...... 165

5.13 Comparison of the intrinsic viscosity of xanthan solutions before and after ultrafiltration ...... 167

xvi 6.1 Schematic diagram of ultraflltration system ...... 176

6.2 Schematic diagrams of the experimental setup for pressure drop measurement ...... 177

6.3 Determination of the empirical equation for UF process ...... 180

6.4 Determination of K and n ...... 182

6.5 Effect of xanthan concentration on K and n ...... 183

6.6 Effect of xanthan concentration on pressure drop in UF module ...... 185

6.7 Effect of pumping rate on pressure drop in UF module ...... 186

6.8 Effect of xanthan concentration on power consumption, filtrate rate, and power consumption rate...... 188

6.9 Effect of pumping rate on power consumption, filtrate rate, and power consumption rate ...... 190

6.10 Effect of transmembrane pressure on power consumption, filtrate rate, and power consumption rate...... 191

6.11 Operating costs for UF at transmembrane pressure equals 70 psig under various shear rate with: (a) membrane cost at $200/m2; and (b) membrane cost at $40G/m2...... 193

6.12 Operating costs for UF at shear rate equals 153 s*’ under various transmembrane pressure with membrane cost at $200/m2 ...... 194

6.13 Total energy consumption for xanthan gum recovery process using UF followed by alcohol precipitation ...... 196

A.l Sample analysis procedures ...... 219

A.2 Calibration curves for cell density and plate count number versus OD ...... 220

A.3 Xanthan gum concentration determination ...... 223

xvii ABSTRACT

Xanthan gum fermentation using immobilized cells of Xanthomonas campestris in

a centrifugal, packed-bcd bioreactor (CPBR) was studied- Cells were immobilized in a

rotating fibrous matrix and the bioreactor was operated as a repeated batch reactor, where

the immobilized cells were repeatedly used for xanthan production. Cell-free xanthan

broth was produced, with -85% product yield from glucose. The xanthan productivity

and the quality of xanthan remained unchanged in every subsequent batch cycle for the

entire period of over 3 weeks studied. The volumetric xanthan productivity in CPBR was

~1 g/L h, which was 2-3 times higher than that from conventional batch fermentation in

the stirred tank reactor (STR) with free cells. The higher productivity was attributed to

the higher cell density, -7 g/L, attained in the CPBR. However, the specific xanthan

productivity was lower in CPBR than in STR. This was because the relatively low cell

viability in CPBR, which might have been caused by oxygen limitation.

The volumetric mass transfer coefficient, kia, and oxygen transfer in CPBR was studied and compared with other bioreactor systems, including STR with disc turbine

(STR-DT), STR with marine propeller (STR-MP), and STR with water-in-oil emulsion

(STR-WIO). It was found that at xanthan concentrations lower than 2%, the kia values in

CPBR were slightly lower than those of STR-WIO and STR-DT. However, at xanthan concentrations higher than 2%, CPBR had higher kia than that of STR-DT and almost the same kLa as that of STR-WIO. This indicated that CPBR had improved oxygen transfer rates at high xanthan concentrations. Ultraflltration as a method to concentrate dilute xanthan fermentation broth before xanthan recovery by alcohol precipitation was studied. A polysulfone membrane hollow fiber (with 500,000 MWCO) tubular cartridge was used. The highly viscous xanthan fermentation broth at its normal concentration of -2.5 (w/v) % was concentrated to 13.5

(w/v) % or higher, with a recovery yield of -95% or higher. The process did not give any adverse effects on the xanthan polymer. The ultrafiltration process was relatively stable and no significant fouling occurred. The filtrate flux (Jv)decreased with increasing xanthan concentration (C) because of the increased solution viscosity. It increased with increasing pumping (Q) and shear (y) rate mainly because of the shear-thinning property of xanthan solution. Filtrate flux also increased with increasing trans-membrane pressure difference (APm). An engineering equation was developed to model Jv as functions of C,

APm, and y. The major power and energy consumptions in ultrafiltration are due to high- rate pumping. The pressure drop for pseudoplastic fluid (xanthan solution) flow in lube was studied. An engineering equation was developed to model the process. Total energy consumption and operating costs for xanthan ultrafiltration were estimated and found to be significantly lower than those with the current alcohol precipitation process. Thus, ultrafiltration can be used economically to concentrate xanthan broth from fermentation by a factor of five or higher, thereby reducing the amount of alcohol needed for xanthan recovery by at least 80%. Energy costs are also reduced by 80%.

xix CHAPTER /

INTRODUCTION

LI BACKGROUND

1.1.1 Xanthan Gum Fermentation

Xanthan gum is extensively used in food, cosmetic, pharmaceutical, and oil-

recovery industries as a thickening, stabilizing, and suspending agent (Cottrell et al .,

1980). Xanthan gum is a microbial polysaccharide commercially produced by fermentation of glucose with Xanthomonas campestris (Margaritis and Pace, 1985). The worldwide consumption of xanthan gum is estimated at over 50 million lbs per year, with an annual growth rate of 1%,

The present industrial process for xanthan gum production is energy-intensive and costly (Figure 1.1). This is mainly because the high viscosity of xanthan solution causes the agitation and aeration in the conventional stirred tank fermentor to be extremely difficult and consequently limits the final xanthan concentration from fermentation to below -?>%

(wt/v) and xanthan productivity to -0.5 g/Lh. The recovery of xanthan gum from fermentation broth is generally done with alcohol precipitation. The large amount of alcohol (isopropanol or ethanol) required in the subsequent recovery process is costly, even

1 Alcohol Recycle f Fermentor Alcohol Laboratory Precipitation □ and Washing Distillation 3#/0 d , U-A alcohol c o Heat i I Drying Sterile Air T I Waste Stream Xanthan Gum

FIGURE 1.1: Conventional process for xanthan gum production. 3 with nearly complete recovery of the alcohol by distillation (Gonzales etal., 1989). The distillation process for alcohol recovery is not only energy-intensive but also causes air pollution. Therefore, an energy efficient, environmentally benign, and cost effective process for xanthan gum production is needed.

There have been many attempts to increase xanthan concentration and xanthan production rate and to lower the energy costs for viscous xanthan fermentation by using improved agitation designs (Funahashi et a i, 1987; Galindo and Nienow, 1992; Galindo and Nienow, 1993; Herbst etal ., 1987; Herbst etal ., 1992; Himmelabach, 1985; Nakajima et a t, 1990; Nienow, 1984; Peters etal ., 1992; Tecante and Choplin, 1993), new types of bioreactors (Kawase and Tsujimura, 1994; Kessler et a i, 1993; Misra and Barnett, 1987;

Pons et a i, 1989; Suh et al., 1991; Zaidi et al ., 1991), and new fermentation systems such as water-in-oil emulsion that has improved aeration and oxygen transfer (Ju and Zhao,

1993; Robinson and Wang, 1988; Schumpe et al., 1991). Some of these fermentation studies have reported a final xanihan concentration of 5% or higher, but with a significantly lowered productivity. Most of these fermentation improvements cannot be used to economically produce xanthan gum at a large scale. A new xanthan production process that is energy efficient and cost effective is thus needed. Other recent research efforts thus have focused on either improving alcohol precipitation (Gonzales et al., 1989; Flahive et al.,

1994) or developing new post-fermentation concentration and recovery methods, such as the ultrafiltration process studied in this work.

1.1.2 Immobilized Cell Bioreactor

Immobilized cell reactors have been used successfully for many years in wastewater treatment and in the vinegar industry (Atkinson et a i, 1984). Cell concentration in an 4 immobilized cell reactor can be severalfold higher than that in free cell systems (DiCosmo et at ., 1994). This trait may compensate for any reduction in cell activity caused by diffusional limitations (Black, 1986). The ability to operate without cell growth is an important characteristics of immobilized cell reactors and allows the possibility for continuous production of secondary metabolites with more efficient bioconversions. It has been demonstrated that by choosing a glucose-based medium that will support product formation but not cell growth, the fermentation reaction can be diverted to produce desirable products instead of simply increasing their biomass (Foerberg et at., 1983;

Mosbach et a i, 1983; Yang et ai ., 1994).

Observed benefits of immobilized cell systems are high reactor productivity at high throughput, ease of biomass recovery, and enhanced gas-liquid mass transfer in the system

(Atkinson, 1986). Cell immobilization also favors cell separation, allows fermentors to be operated on a drain and fill basis, and greatly facilitates recycling or reuse of microorganisms.

The idea of using cell immobilization technology to improve xanthan fermentation is not new. Xanthan gum fermentation using cells immobilized in porous celitc beads to control broth viscosity in the fermentor was studied by Robinson and Wang (1988). In spite of the limited oxygen transfer rate inside carrier particles, simultaneous production and concentration of xanthan gum, reaching -5% xanthan in the particles, was achieved.

However, separating and recovering xanthan gum form the particles was very difficult.

A fihrous-bcd immobilized-cell bioreactor developed by Yang et at. (1994) has been demonstrated to provide natural attachment for reaching high cell density and facilitate the separation of cells from products in solutions. Growth of cells to a high density of 30 to

100 g dry weight per liter reactor packed volume can be achieved in this fibrous bed bioreactor. The structured fibrous bed allows for good multiphase (gas, liquid, solid) 5 flows and provides renewable surfaces for cell immobilization. This reactor is also able to operate continuously for a long period without encountering any clogging or degeneration problems, which often occur in conventional immobilized cell and membrane bioreactors.

This fibrous bed bioreactor has been successfully used in fermentations to produce propionic acid (Yang et at., 1994). This immobilized cell bioreactor provides a new approach towards the production of fermentation products. Cell immobilization with fibrous matrix also provides an opportunity to produce cell-free xanthan broth, which is important to the applications in tertiary oil recovery.

The high viscosity of xanthan solution at low concentration presents a major challenge in mixing xanthan broth during fermentation (Zaidi et al., 1991). The agitation and aeration (oxygen transfer) in a conventional stirred-tank bioreactor are not appropriate for the immobilized cell bioreactor. However, the xanthan solution shows a high degree of pseudoplasticity, i.e., the viscosity decreases rapidly as the shear rate increases (Kang and

Pcttitt, 1993). This shear-thinning property allows efficient pumping of xanthan polymer at high pumping (shear) rates. The mixing problem thus may be overcome by continuous medium circulation through the fibrous matrix and by rotating the fibrous matrix to enhance air, liquid, and cell contacts and to separate the xanthan polymer from the cells. In this new design, liquid media and air is passed through the porous fibrous matrix to ensure intimate contact with the immobilized cells. The centrifugal force generated form rotating the device is high enough to separate xanthan polymer from cells. The xanthan gum production rate in the immobilized cell bioreactor will be dependent on the cell density and the oxygen transfer rate (OTR) in the fibrous matrix. It is know that at a high rotational speed or under high gravitational force (> 2-3 g) conditions, gas and liquid can be passed through a contacting medium at high rates without flooding to achieve a high mass transfer rate that is

2 or 3 orders of magnitude higher than that obtained in conventional mass transfer equipment (Mohr and Khan, 1987). Thus, when operated at a high rpm, the centrifugal 6 fibrous bed bioreactor should be able to provide a much higher OTR for the viscous xanthan broth and will greatly improved xanthan production rate.

1.1.3 Ultrafiltration of Xanthan Solution

Fermentation broths are complex mixtures of biomass, dissolved macromolecules, and electrolytes (Fane and Radovich, 1990). Separation and concentration of these materials are difficult because the desired products are usually in dilute solution and have similar physical and chemical properties. The recovery of xanthan gum form fermentation broth is usually done with alcohol precipitation, which is expensive and accounts for more than 50% of the total xanthan production costs. The final xanthan product from alcohol precipitation is not pure, often containing cells, cell debris, proteins, and salts as co­ precipitates. Alternative recovery methods, including direct drum-drying, spray-drying, and freezing-thawing, have been proposed to eliminate the use of alcohol (Rogovin e ta i ,

1965; Haze et o/., 1989). However, the energy demands for these processes are exceedingly high.

Ultrafiltration offers the advantage of requiring low pressures for operation and having relatively low equipment and operating costs as compared to other membrane processes. Membrane-based ultrafiltration serves as a molecular sieve to separate solute molecules of different size and provide a means of separation and concentration al the molecular and fine-particle level (Cheryan, 1986; Barker and Till, 1992). In ultrafiltration, liquid is pumped across the membrane surface at high velocity to prevent cake formation

(Shiloach and Kaufman, 1986; van Oers et al ., 1992), resulting in high filtration rates that can be maintained continuously (Lasky and Grant, 1985). Although ultrafiltration has been widely studied and used in industry for protein concentration and waste stream clean-up 7 (Kimura, 1992), no report can be found for its application to concentrating highly viscous

polysaccharide solutions.

The very high viscosity of xanthan solution at low concentrations presents a major challenge in using ultrafiltration to concentrate xanthan broth, as the filtrate flux generally

decreases rapidly with increasing the solution viscosity. However, the xanthan gum

solution shows a high degree of pseudoplasticity, i.e., the viscosity decreases rapidly as

the shear rate increases (Kang and Pettitt, 1993). This shear-thinning property allows efficient ultrafiltration of xanthan polymer at high pumping (shear) rates. It is thus possible

to use the ultrafiltration process to remove water and impurities present in the xanthan

fermentation broth to produce a high-concentration, purified xanthan solution.

1.2 OBJECTIVES

The goal of this research was to develop a novel, energy-efficient process consisting of an immobilized cell bioreactor for xanthan gum fermentation and an ultrafiltration process to concentrate xanthan fermentation broth (Figure 1.2). In this new process, the cell-free xanthan broth produced in the immobilized cell bioreactor is concentrated by ultrafiltration to 15% (wt/v). This process thus can reduce the amount of alcohol used in xanthan recovery by at least 80%. Water (and other medium components) used in fermentation can also be recovered from ultrafiltration and recycled for fermentation use, further reducing wastes generated from this process. Therefore, the objectives of this work included: Fermentation Alcohol Recycle Stage I Stage II Growth Production Centrifugation Ultrafiltration or Filtration Distillation 15 %

alcohol Heat Alcohol Precipitation and Washing Filtrate T Drying Waste Stream Feed Glucose and Nutrients I Xanthan Gum

FIGURE 1.2: Proposed process for xanthan gum production. 9 1. To develop a novel centrifugal, packed-bed bioreactor for xanthan fermentation:

A novel, centrifugal, packed-bed bioreactor modified from the columnar bioreactor

with spirally wound fibrous bed was employed for xanthan gum production. Cell

immobilization technology was employed to reach high cell density, whereas two-

stage repeated batch process was used to perform high xanthan production rate with

the aid of enhanced oxygen transfer rate by applying centrifugal force and medium

recirculation. Several key factors affecting xanthan gum production were identified

in order to optimize the xanthan fermentation. The aeration capability of the

centrifugal, packed-bed bioreactor was also evaluated in terms of its oxygen

transfer rate in comparison with other bioreactor systems.

2. To evaluate the ultrafiltration process for concentrating xanthan broth:

The feasibility and capability of ultrafiltration process for xanthan broth

concentration were studied, as well as the membrane phenomenon encountered

during xanthan ultrafiltration at various conditions. The properties of xanthan gum

product obtained from ultrafiltration processes were evaluated by means of

rheology evaluation and molecular weight determination. The energy consumption

of the ultrafiltration process was estimated in order to compare the operating costs

for this process with ethanol precipitation and distillation.

1.3 SCOPE OF THE STUDY

The goal of this research was to develop a new process to produce xanthan gum at a significantly lower cost than that of the present industrial process. In order to completely understand the properties and process requirements for xanthan gum production, a 10 literature review of xanthan gum and ultrafiltration is given in Chapter II. From Chapter III

to Chapter VI the format of each chapter as an independent article for journal publication

was used.

Chapter HI includes the design and development of the new centrifugal, packed-bed

reactor (CPBR), In order to shorten the fermentation time required for conventional

xanthan gum production, major improvements that could be characterized into four

categories were employed in the CPBR, including cell immobilization, rotational force,

circulation stream, and two-stage operation (Figure 1.3). The cells of X. campestris were

immobilized in a rotating Fibrous matrix inside the bioreactor. The centrifugal force from

rotating the fibrous matrix separated xanthan polymer from the immobilized cells in the

matrix. The cells immobilized in the bioreactor were repeatedly used for xanthan

fermentation to achieve the production of cell-free xanthan broth. The liquid medium and

air were passed through the porous fibrous matrix to achieve intimate air, liquid, and cell contact and an enhanced oxygen transfer rate during the fermentation. Different optimal

medium compositions, one for cell growth and one for xanthan production, were used in

first and second stages, respectively, of the two-stage, immobilized cell fermentations. The attachment and morphology of immobilized cells on the fibrous matrix were examined with scanning electron microscopy, and the viability of the immobilized cells was also investigated by plate count method.

In Chapter IV, the performance of xanthan fermentations in the CPBR with immobilized X. campestris was evaluated in terms of oxygen transfer capability and compared with different systems, including a stirred tank bioreactor with disc turbine

(STR-DT), STR with marine propeller (STR-MP), and STR with water-in-oil emulsion

(STR-WIO). Effects of the circulation stream with a nozzle spray system on oxygen IMMOBILIZATION RECIRCULATION

1. Hydrophilic matrix 1. Provide gas-liquid­ 2. High cell density cell contact 3. Cell Viability 2. Spray medium on 4. Reuse of cells matrix 5. Cell-free broth 3. Enhance mixing

BIOREACTOR

ROTATION TWO STAGES

1. Generate centrifugal 1. Temperature force to remove 2. pH xanthan gum 3. Oxygen 2. Provide agitation 4. C/N in medium

FIGURE 1.3: Characteristics of the centrifugal, packed-bed bioreactor (CPBR). 12 transfer in CPBR at various xanthan concentrations are also studied to determine the importance of circulation stream in fermentation aeration.

Chapter V discusses the use of membrane ultrafiltration to concentrate xanthan gum in the fermentation broth. The feasibility and kinetics of using a hollow fiber ultrafiltration unit for concentrating dilute xanthan broth produced from fermentation from -2.5% to

-15% (wt/v) were studied. Several process parameters, including xanthan concentration, transmembrane pressure difference, pumping (shear) rate, and solution pH, were studied to evaluate their effects on filtration performance. Also, possible shear effects from ultrafiltration on the xanthan polymer were investigated by examining the Theological properties and molecular weight of xanthan gum before and after ultrafiltration process.

Chapter VI gives the estimated operating costs of the ultrafiltration process based on its energy consumption and membrane area required. In order to estimate the energy consumption in the ultrafiltration process, the pressure drop that depends on the flow behavior of xanthan gum solutions in the ultrafiltration module was studied using the power law and momentum balance equations. The actual pressure drops inside ultrafiltration module under various xanthan concentrations were measured in comparison with the theoretical values. The operating costs of ultrafiltration were then compared with the energy costs of distillation in recovering the alcohol used for xanthan precipitation.

Chapter VII concludes this study with an overview discussion of its results.

Conclusions about the success of the project are made and discussed, and recommendations for future work are presented. CHAPTER II

UTERA TORE REVIEW

2.1 INTRODUCTION

Xanthan gum is a branched, anionic extracellular heteropolysaccharide with a

reported molecular weight ranging from 2 million to 13-50 million (Dintzis et al., 1970;

Holzwarth, 1978). Xanthan gum is commercially produced by fermentation of glucose

using Xanthomonas campestris, a rod-shaped, gram-negative, strictly aerobic bacterium

which also causes vascular disease of cabbages, cauliflower, and rutabagas (Lilly et al.,

1958; Kennedy and Bradshaw, 1984).

Among all Xanthomonas campestris, NRRL B-1459 produces the most significant amount of stable viscous water-soluble xanthan polysaccharide in fermentation. It was first screened by the Northern Regional Research Laboratory (NRRL) of the U.S. Department of Agriculture during the 1950's. As the first microbial gum commercially available by large scale fermentation since 1964 (Kang and Cottrell, 1979; Cnieger and Crueger, 1984), xanthan gum has become well established in both food and industrial applications as a thickening, suspending, stabilizing, and gelling agent because of its excellent and unique rheological properties (Margaritis and Pace, 1985). Since 1969, xanthan gum has been classified as a GRAS (generally recognized as safe) food additive by The Food and Drug

13 14

Administration (FDA), and this made it a more widely used product than other microbial

gums and synthetic polymers in food, pharmaceutical, and chemical industries (Zatz,

1984). Xanthan gum is also used as a mobility control agent in chemically enhanced oil

recovery (CEOR), where it is used to increase the viscosity of drive waters used to force oil

trapped in underground reservoirs to the surface by polymer flooding (McNeil and Harvey,

1993).

Highly viscous xanthan solutions can be formed from the cellulose-like backbone

of the molecule, even at a low concentration in either hot or cold water (Kennedy and

Bradshaw, 1984). Xanthan can also form elastic, thermoreversible gels when combined

with locus bean gum, and can achieve extremely high viscosities when combined with guar

gum. Xanthan gum solutions exhibit high stability at extremes of pH and shear, and high

resistance to thermal degradation (McNeely and Kang, 1973). The stability and viscosity

of the resulting xanthan solution are claimed to be enhanced by the addition of salts such as

potassium chloride, sodium chloride, or calcium chloride (Patton, 1976; Cottrell et al .,

1980). However, to increase xanthan utilization in expected commercial areas, a significant

increase in the present production level and reduction in its production costs are required.

In this chapter, brief reviews of chemical structure and rheological properties of xanthan

gum, xanthan gum fermentation and recovery processes, and ultrafiltration are provided.

2.2 STRUCTURE AND CONFORMATION OF XANTHAN

The basic repeating unit of xanthan macro-molecule was first proposed to be composed of D-glucose, D-mannose, D-glucuronic acid, acetic acid, and pyruvic acid by

Sloneker and Jeanes (1962). The molar ratios of D-glucose, D-mannose, and D-glucuronic acid in xanthan were then found to be approximately 2:2:1 (Sutherland, 1977). The 15 currently accepted structure of xanthan was initially proposed by Jansson etal. in 1975.

As shown in Figure 2.1, the structure consists of a linear backbone of p-(l-4) -linked D- glucose residues which has three-unit-long side chains appended on alternate residues on the main chain at the C-3 position. D-Mannose residues directly appended to the backbone bear D-acetyl substituents on the C -6 position. Pyruvic acetal substituents, i.e., 4,6-0-(l- carboxyethylidene), are on the terminal D-mannosyl residues of some of these side chains; their frequency of occurrence depends on the bacterial strain and fermentation conditions.

The anionic acetyl and pyruvic acetal groups also may contain monovalent cations at neutral pH (Cottrell and Kang, 1978; Sanderson, 1981).

Different conformational forms of xanthan gum in aqueous solution have been investigated by electron microscopy. As shown in Figure 2.2, xanthan polymeric chains can exist as single-stranded, double-stranded, or partly dissociated double-stranded structures (Stokke et a l, 1986). The quantity of pyruvate in xanthan polymer varies from virtually 0% to its theoretical maximum of 7.5%, and the amount of acetate is also variable

(Jansson e ta l, 1975; Cadmus e ta l, 1976; Davidson, 1978; Sutherland and Ellwood,

1979; Jarman and Pace, 1984; Sutherland, 1982).

Recently, new research efforts were made to obtain additional molecular information on xanthan polymers, including the location of a second O-acetyl group

(Stankowski e ta l, 1993), the structure of xanthan in terms of O-specificity (Bukharov et a l, 1993), the image data from scanning tunneling microscopy and transmission electron microscopy (Wilkins e ta l, 1993), and the radiowave dielectric properties of xanthan in aqueous solutions (Bordi e ta l, 1995). CK.OH

OH

M = Na, K, 1 /2 Ca

OH

FIGURE 2.1: Repeated unit of xanthan gum molecules. Disordered or random Ordered or rigid Aggregate coil subunit rod subunit AA mm

Favored bv Favored bv Favored bv

Low concentration and High concentration and high ionic strength high ionic strength Low pyruvate High pyruvate Low acetate to pyruvate ratio High acetate to pyruvate ratio High temperature Low temperature Low ionic strength High ionic strength

FIGURE 2.2: Conformational forms of xanthan in solution. 18 2.3 CHARACTERISTIC PROPERTIES OF XANTHAN GUM

2.3.1 Pseudoplasticity

Rheology is defined as the science of the deformation and flow of matter (Brodkey and Hershey, 1988). The rheology of xanthan solutions is strongly influenced by molecular association and aggregation. Xanthan gum in aqueous solution is highly non-

Newtonian in character - the viscosity is a strong function of the shear rate. This is found to be important in determining the mixing, heat transfer, and oxygen transfer requirements of xanthan fermentation (Lim et al ., 1984).

Water solutions of xanthan gum at low concentration show pseudoplasticity, an unusual Theological property as compared to other thickeners. The apparent viscosity of the pseudoplastic solution decreases with increasing shear rate (Jeanes, 1974; Pace and

Righelato, 1980; Cottrell etal.. 1980). A basic shear diagram illustrating the various types of non-Newtonian behavior of aqueous solutions can be obtained by plotting the logarithm of shear stress (x) versus the logarithm of shear rate ( - y ) (Ostwald, 1924). If the plot is linear with slope unity, then Newton's law is obeyed. When a complete basic shear diagram is determined, the slope of a pseudoplastic liquid is less than one (Brodkey and

Hershey, 1988). The Ostwald-deWaele equation (Ostwald and Auerbach, 1926), commonly called the power law, is often used to correlate the shear stress and the shear rate:

r = * ( - y ) fl (2.1) where ris the shear stress (Pa, or dyn/cm2), - y is the shear rate (s 1). « is the flow behavior index (-), and K is the consistency index (Pa s"). 19 The xanthan solution also may exhibit a yield stress, a given stress which has to be applied to the fluid before movement occurs. The yield stress is related to the structure of xanthan molecules in solution (Jeanes, 1974; Whitcomb and Macosko, 1978; Solomon et at., 1981; Lim et at., 1984; Honnote et al., 1991). With the yield stress ( ro), equation

(2 .1) can be modified to;

T=zo+K(-yy (2.2)

Xanthan solution rheology changes with ionic strength, pH, and temperature.

Xanthan solutions at low ionic strength undergo a thermal transition with the viscosity change (Morris etal., 1977). Generally, uniform and high viscosity of xanthan solutions can be maintained over the pH range 2-12, with some reduction at extreme pH values

(Kang and Burnett, 1977). Xanthan solutions are able to retain their viscosities up to around 80°C (Symes, 1980). The Theological properties of xanthan macromolecules appear to be primarily affected by the content of pyruvate residues, which promote the aggregated form of xanthan (Smith and Pace, 1982; Smith et al., 1984; Margaritis and

Pace, 1985). The molecular weight of the polymer also has a direct influence on its

Theological behavior, with larger molecules having higher viscosity. Although it is known that the molecular weight of xanthan can vary, little is known about the termination of polymer chain formation and how the growth medium can be altered to control it

(Margaritis and Pace, 1985). The dissolved oxygen concentration provided during xanthan fermentation appears to be an important factor to the xanthan molecular weight produced from the fermentation (Flores et al., 1994).

The flow behavior of low-concentration (up to -2%) xanthan broth was studied by

Galindo et al. (1989). The xanthan rheology was found to be highly influenced by gum concentration, and the differences in flow behavior were associated with different molecular conformations of xanthan polymer (Lim et al., 1984; Tam and Tiu, 1993; Ovialt 20 and Brant, 1994; Rochefort and Middleman, 1987; Garcfa-Ochoa and Casas, 1994). In general, the Theological parameters were found to vary in the following ranges: K (2-210 Pa s'1), n (0.2-0.9), and xo (0.2-50 Pa).

2.3.2 Molecular Weight

A wide range of values from 2 to 50 million dalton has been reported for the molecular weight of xanthan polymer as determined by either classical light-scattering techniques or band-sedimentation analysis (Dintzis etal., 1970; Rinaudo and Milas, 1978;

Holzwarth, 1978; Kennedy and Bradshaw, 1984). It was recently reported that the mean molecular weight of xanthan equaled 1 0 million dalton when the dissolved oxygen tension

(DOT) was maintained above 40% in the medium during fermentation (Flores et al., 1994).

Moreover, it was also claimed that, in general, the proportion of higher molecular weight polymer increased with increasing DOT.

The intrinsic viscosity of polymers provides a useful starting point for rationalizing the viscosity of flexible polymer chains and the molecular weight of a polymer (Eisenbcrg and Crothers, 1979). The addition of a polymer to a solution causes the viscosity to increase. The reason is that the polymer extends across the flow planes and increases the frictional resistance to shear. The increment in the viscosity is very sensitive to both polymer size and shape. The viscosity increment is defined in terms of the specific viscosity, tj sp;

where t j is the polymer solution viscosity and t j 0 is the viscosity of the solvent without polymer. Milas et al. (1985) concluded that the specific viscosity at zero shear rale is a 21 function of the molecular weight of xanthan and the polymer concentration in its ordered conformation.

The viscosity of a polymer solution is a complex function of the concentration

because, at high concentration, the distortion of the velocity pattern of the liquid by one molecule can affect the shear at a neighboring molecule. Thus, in general, one can expect

that the viscosity of a solution should be expressed as a power series relative to the

viscosity 77 0 of pure solvent:

77= 770 (l + V + * 2cJ+-'-) (2.4)

The relative viscosity, 77 re), is 77/77 0, and from equation (2.2) we can see that

V sp — V rel * 1 ~ kjC + k2C^ + (2.5) so 77 sp is the increase in the relative viscosity of the solution beyond the value 1 for pure solvent.

For low concentrations of added polymer, the viscosity increment, 77 sp, is proportional to the concentration. The limiting ratio can be defined as [ 77]

in which [ 77I is called the intrinsic viscosity, and c is the polymer concentration in weight/unit volume (Torres et a t, 1993). Intrinsic viscosity has the unit of cm 3 g_l

(Eisenberg and Crothers, 1979).

It is [77] that can be computed by considering the extra energy dissipation that a suspended spherical molecule of molecular weight M causes in a fluid under shear. The 22 molecular weight of xanthan gum can be further evaluated by use of the Mandelkem-FI ory-

Scheraga (MFS) equation (Holzwarth, 1978; Cantor and Schimmel, 1980):

(2.7)

where s is the sedimentation coefficient, rj 0 is the solvent viscosity, v is the partial specific volume of the polymer, p is the solvent density, Na is the Avogadro's number, and fi has a value near 2.5 x 10 6 when [ 17] has the units dl/g (Mandelkem and Flory, 1952;

Scheraga and Mandelkem, 1953).

The sedimentation coefficient, s, is also found to be a function of molecular weight,

M(Dintzis etal., 1970). According to Holzwarth (1978), the intrinsic viscosity [T 7J obeyed the relation [tj] = KM135. The value of M obtained via the MFS equation (equation (2.7)) for commercial "Keltrol" (Kelco, CA) native xanthan was 12.2 x 10 6 with [ 77] at 12,3(H) mL/g, whereas a xanthan culture broth grown in Holzwarth's laboratory showed 52 x 10 6 in M with the value of 35,000 mlVg in [tj].

2.4 MICROBIAL BIOSYNTHESIS OF XANTHAN GUM

2.4.1 Function and Importance

Margaritis and Pace (1985) pointed out the importance of understanding the pathways involved in exopolysaccharide biosynthesis, because this information can he used to control and optimize rates and yields of fermentation, as well as the physicochemical characteristics of the resulting exopolysaccharides. 23 The production of extracellular polysaccharide was proposed to have a survival function for the bacteria, protecting microorganisms against various adverse environmental conditions. For Xanthomonas campestris, in particular, the high moisture-retaining capacity of xanthan gum provides the organism with a protective layer under low humidity environmental conditions; the slime layer, xanthan, may act as a barrier to protect the cell against attack from bacteriophages; and xanthan may contribute to the recognition of appropriate sites on the host plant at the initial stage of colonization of the bacteria

(Kennedy & Bradshaw, 1984).

2.4.2 Enzyme System

Exopolysaccharide synthesis is suggested to be a multienzymatic system

(Sutherland, 1977), and each enzyme assumes a high degree of specificity (Kang and

Cottrell, 1979). The enzymes involved in xanthan synthesis may be classified into four groups (Sutherland, 1977): (i) Group I: those enzymes which are involved in the initial metabolism of the substrate, such as hexokinases; (ii) Group II: enzymes responsible for the synthesis and interconversion of sugar nucleotides, such as UDP-glucosc pyrophosphorylase and UDP-glucose dehydrogenase; (iii) Group III: transferases which are responsible for the formation of the repeating monosaccharide unit attached to the carrier lipid; and (iv) Group IV: translocases or polymerases which form the xanthan molecule. 24 2.4.3 Mechanism of Biosynthesis

Four steps of reactions are considered to be responsible for the synthesis of xanthan gum during fermentation processes.

The First step is the uptake of the substrate (carbohydrate), which may occur by active transport, group translocation, or facilitated diffusion. Group I enzymes are involved in this step, and the substrates are phosphorylated.

The second step in the biosynthesis is the intermediary metabolism, involving the formation of precursors for assembly of the xanthan repeating unit. Group II enzymes activate monosaccharides into immediate precursors, such as UDP-glucose, GDP- mannose, and UDP-glucuronic acid.

The third step occurs at the cell periplasmic membrane. Group III enzymes transfer the sugar nucleotides to isoprenoid alcohol phosphate (a lipid acceptor) to form a pentasaccharide-P-P-lipid. The pyruvic acid acetal residues in xanthan are formed by transfer from phosphenolpyruvate (PEP) to the pentasaccharide-P-P-lipid (Ielpi et al .,

1981), The addition of acetyl residues also occurs at the lipid intermediate stage and the donor is acetyl-CoA (Kennedy and Bradshaw, 1984).

The fourth step is polymerization which occurs by transfer of the pentasaccharide unit from the lipid carrier to the growing xanthan (Kennedy and Bradshaw, 1984). Figure

2.3 shows a generalized xanthan biosynthesis pathway proposed by Jarman and Pace

(1984). Glucose Fructose 6-P i — ATP ♦ Glucose 6-P Mannose 6-P Glucose 1-P f— NADPH2 Mannose 1-P LTTP 6-Phosphogluconic acid UMP IP GTP 1 nnrDugol pathway Entner-Doujgrofl GDP-Mannose IDP-Glucose UDP-Glucose 2-Oxo-3-deoxy-fFphosphogluconic acid GDP J ______2NAD(P)H2 r via IDP-linked intermediates UDP-Glucuronic acid Glyceraldehyde 3-P i h— ^ NADHj (Glu-Glu)n 1,3-oiP-Glyceric acid | ^ ATP Mann-6-O-Ac Pholphog lycerate GluA (Na-*) ATP t nolpyruvate Pyruvate Mann {4,6 Pyr NADH Acetyl Co A

(IP = Isoprenoid phosphate, IDP = Isoprenoid pyrophosphate) Tricarboxylic acidJ cycle

FIGURE 2.3: Proposed pathway of xanthan synthesis in Xanthomonas campestris (adapted from Jarman and Pace, 1984).

ro 26 2.5 FERMENTATION KINETICS

Xanthan fermentation kinetics has been extensively studied (Moraine and Rogovin,

1973; Weiss and Ollis, 1980; Pinches and Pallent, 1986). During batch fermentation, cells undergo a physiological change from trophophase to idiophase. After entering idiophase the cells stop growing, but xanthan gum continues to accumulate (Behrens et al., 1980).

Typical data on cell growth, xanthan production, and glucose consumption with pH controlled at 7 for a xanthan fermentation at 30°C are shown in Figure 2.4. At the early stage of fermentation, the inoculated culture showed exponential growth. The culture then reached the stationary phase. Xanthan gum concentration increased with time until all glucose was consumed. Dissolved oxygen tension (DOT) monitored in the medium decreased as the fermentation progressed, dropping gradually from 1 0 0 % initially to approximately 20%. As the broth became gel-like and viscous, the DOT dropped to approximately 10%. This phenomenon mainly depended on the oxygen consumption for both cell growth and xanthan production. The oxygen transfer rate (OTR) in the medium was reduced owing to increased broth viscosity resulting from increased xanthan concentration.

2.5.1 Cell Growth Kinetics

In a typical batch process the number of living cells varies with time. After a lag phase, where no increase in cell number is evident, a period of rapid growth ensues, during which the cell number increases exponentially with time. This stage of batch culture is referred to as the log, or exponential, growth phase (Bailey and Ollis, 1986).

The growth rate of cells in the exponential phase is first order, and it may be described by the following equation: 27

30

Cell

I 20 - Glucose e °8 0 C •** □ es Xanthan e V w c 10

o Cell Density (g/L) U

20 4060 80 100 80 £ 60 oH 40 Q 20 0 20 60 6040 0.3 £ 0.2 0.1 H o 0.0 # 0 20 4060 60 600 5 « 500 o ? 400 o U 300 IS s w 200 100 CWa w S u 0 20 40 60 80 Time (h)

FIGURE 2.4: A typical batch fermentation for xanthan gum production by Xanthomonas campestris (30°C). 28

(2.8) where X is the concentration of cell biomass, t is time, in hour, and p is the specific growth rate , in hour1.

On integration, equation (2.8) gives:

(2.9) where X0 is the original biomass concentration and Xt is the biomass concentration at time t. On taking natural logarithms, equation (2.9) becomes:

In X, = In X0 + p t (2. 10)

Thus, a plot of the natural logarithm of biomass concentration against time should yield a straight line, with its slope equals to fl

A linear relationship between the specific oxygen uptake rate (OUR, in the unit of mg O2 / g cell /h) and the specific growth rate, ft, was found by Pinches and Pattern

(1986). The OUR increases with increasing dissolved oxygen level up to a critical point above which no further increase in OUR occurs. Thus, maximum biomass production may be achieved by satisfying the organism's maximum specific oxygen demand by maintaining the dissolved oxygen concentration greater than the critical level. If the dissolved oxygen concentration falls below the critical level, then the cell may be metabolically disturbed

(Stanburg and Whitaker, 1984).

According to Lo (1993), the value of p for X. campestris NRRL B-1459 during fermentation ranges from 0.06 to 0.12 h r1 depending on the fermentation conditions. 29 2.5.2 Product Formation Kinetics

Ollis and co-workers (Weiss and Ollis, 1980; Klimek and Ollis, 1980) found that the Leudeking-Piret equation can be used to describe the product formation for xanthan gum in batch cultures of X. campestris. The equation is given as;

dP dX — = m ------1-nX (2.11) dt dt where dP/df and dX/dt are the rates of product and cell formation, respectively, X is the cell concentration, and n and m are empirical constants for growth-associated and non-growth- associated xanthan fermentation, respectively. The values of m and n may vary with fermentation conditions, such as temperature (Shu and Yang, 1991) and pH (Weiss and

Ollis, 1980).

Substrate consumption behavior requires modification of the Leudeking-Pirct equation with a maintenance term, and the modified model can be expressed as:

=a^- + pX (2. 12) dt dt where a and /J are parameters associated with growth and non-growth substrate consumption rates, respectively, and can be expressed as follows:

a = T1X +f 1P

0 =■£- + *, (2.14) * P where and Yx are product yield coefficient and cell yield coefficient, respectively, and ke is specific maintenance rate (Moraine and Rogovin, 1971). The growth associated parameters m and a were found to be independent of the concentration of the limiting nutrient, nitrogen (Pinches and Pattern, 1986). 30 The specific xanthan production rate at a specific time was determined from the xanthan production rate ( dP/dt , the slope of the plot of product concentration versus time) divided by the cell density ( Xs, in g/L). According to Moraine and Rogovtn (1973), the specific xanthan production rate drops from almost 0.5 g/g cell/h initially to less than 0.1 g/g cell/h as the solution became extremely viscous at the end of a batch fermentation.

Under oxygen limitation, the molecular weight was found to decrease linearly with the

OUR resulting in lower viscosity yield of the product, and the specific xanthan production rate of the bacteria became a linear function of the OUR (Peters et at., 1989, Pons et at.,

1989). Experimental data obtained by Lo (1993) indicated that the specific production rate is inside the range of 0.2 to 0.4 g/h/g cell.

The yield coefficient of either biomass production or product formation of a fermentation is based on the amount of substrate consumption. The cell yield is expressed as Yx/s whereas the product yield is expressed as Yp/$.

(2.15)

(2.16)

Typically, for xanthan fermentation the value of Yx/s *s between 0.03 and 0.18, whereas the value of Yp/s varies from 0.79 to 0.88 (Lo, 1993).

2.5.3 Effects of Environmental Factors

In order to achieve desired rates of synthesis and yields of xanthan gum and to maximize the mass transfer capabilities of the reactor, other environmental parameters in addition to medium composition, including temperature, pH, shear, and dissolved oxygen. 31 must be carefully controlled in xanthan gum fermentation broth. These parameters are also prime determinants of the purity of the product, its chemical composition and molecular weight, which in turn determine its end use performance (Margaritis and Pace, 1985).

2.5.3.1 Medium composition

Both the cell yield and specific growth rate of X. campestris in a stimed tank reactor

(STR) were found to decrease with increasing C/N ratio, or more specifically, the glucose to yeast extract ratio in the fermentation medium (DeVuysl e ta i, 1987; Lo, 1993). On the contrary, the xanthan yield and specific xanthan production rate increase with increasing the glucose to yeast extract ratio in the medium. Therefore, a two-stage glucose-to-yeast extract ratio shift from an initial ratio of 2.5%-to-0.3% for cell growth to a ratio of 5.0%- to-0.3% for xanthan production is preferable for xanthan fermentation (Lo, 1993).

2.5.3.2 Temperature

In general, the optimal temperature for cell growth also is the optimum for product formation (Kang and Cottrell, 1979). However, Shu and Yang (1990) found that the optimal temperature range for X. campestris cell growth is 22 to 24°C, while the optimum for xanthan production is 30 to 33°C.

2.5.3.3 pH

Because of the acidic nature of xanthan, medium pH decreases with xanthan formation during the fermentation, and both cell growth and polymer production stop at about pH 5.5 if the solution pH is not properly controlled (Margaritis and Pace, 1985). A pH optimum of 7 has been reported for the production of xanthan gum by X. campestris

(Moraine and Rogovin, 1973), whereas pH 6 is the optimal value for cell growth (Thonart et aL, 1985). 32 2.5.3.4 Shear

Shear was found to have positive effects on xanthan fermentation (Moraine and

Rogovin, 1973). A shear sensitive slime layer exists around the cell that acts as a diffusion

barrier to nutrients. When a high shear rate is applied, the cell slime is reduced or

eliminated from the cell surface, lowering the restriction of diffusion of nutrients and thus

enhancing the biosynthesis mechanism (Moraine and Rogovin, 1973; Peters e ta i, 1989).

It has been observed that the shear stress was particularly important for xanthan

gum production. The specific xanthan production rate and the specific oxygen uptake rate

(OUR) were influenced by the shear stress. Both of these rates increased with increasing

shear stress in the range of 0-40 Pa, beyond which they were constant (Funahashi et a i,

1989).

2.5.3.5 Dissolved oxygen

Oxygen supply is essential to both cell growth and biosynthesis of xanthan, such as

for the oxidation of reduced pyridine nucleotides and sugar acids (Margaritis and Pace,

1985). According to Moraine and Rogovin (1973), if the DOT is maintained above 20%, oxygen has no effect on xanthan fermentation. The specific oxygen demands are 182 mg

0 2 /g cell/h for cell growth during the growth phase, and 78 mg 0 2 /g cell/h for cell maintenance during the stationary phase, and 70 mg 0 2 /g cell/h for xanthan biosynthesis

(Herbst et a i, 1987). Recently, Flores and coworkers (1994) found that DOT in the culture medium strongly affect the rheological behavior and molecular weight of the xanthan gum produced from fermentation. 33

2.5.3.6 Other factors

Other factors affecting cell growth and xanthan production, including strain selection, culture maintenance, and salts in medium formulation, have also been studied

(Moraine and Rogovin, 1971; Cadmus e ta i, 1976; Davidson, 1978; Souw and Demain,

1979; Sutherland, 1979; Kang and Cottrell, 1979; Kennedy and Bradshaw, 1984; Thonart et a i, 1985). Because of the different optimal conditions for cell growth and xanthan formation, a two-stage fermentation process was proposed to replace the conventional batch process (Silman and Rogovin, 1970; DeVuyst etai ., 1987; Shu and Yang, 1990; Lo,

1993).

2.6 OXYGEN TRANSFER IN XANTHAN FERMENTATION

Xanthan fermentation processes are aerobic and, therefore, require the provision of oxygen. Since it is not possible to provide a microbial culture with all the oxygen it needs for the complete respiration of the glucose or any other carbon source in one addition, a microbial culture must be supplied with oxygen during growth at a rate sufficient to satisfy the organisms' demand. It was reported by Stanburg and Whitaker (1984) that the aeration conditions necessary for the optimum production of a fermentation product may be different from those favoring biomass production.

2.6,1 Determination of Volumetric Mass Transfer Coefficient

The provision of oxygen for a fermentation cannot be estimated simply by overall demand, because the metabolism of the culture is affected by the concentration of dissolved 34 oxygen in the broth. Normally, oxygen is supplied to microbial cultures in the form of air,

the cheapest available source of the gas. The rate of oxygen transfer from air bubble to the

liquid phase may be described by the equation:

^ L = kLa(C * -C L) (2.17) dr

where CL is the concentration of dissolved oxygen in the fermentation broth, in mmole dm' dC 3, t is time in hour, —L is the change in oxygen concentration over a time period, i.e. the dt oxygen transfer rate, in mmoles 0 2 dm-3 h 1' kL is the mass transfer coefficient (cm h*1), a

is the gas/liquid interface area per liquid volume (cm2 cm-3), and C* is the saturated dissolved oxygen concentration or the solubility of 0 2 in the solution, in mmole dm'3.

Table 2.1 lists the solubility of oxygen at 1 atm pure oxygen in water at various temperatures. It is noted that oxygen solubility in water is proportional lo the partial pressure of oxygen in the gas phase and is affected by the electrolyte concentration and solution temperature.

It is extremely difficult to measure kL and a separately in a fermentation; therefore, the two terms are generally combined in one term, kLa , the volumetric mass transfer coefficient, with unit of reciprocal time, normally h*1. The volumetric mass transfer coefficient is a measure of the aeration capacity of a fermentor under the test conditions.

The larger the kLa, the higher the aeration capacity of the system. The determination of kta of a fermentor is essential in order to establish its aeration efficiency and to quantify the effects of operating variables on the provision of oxygen.

Gassing-out technique is one method commonly used to estimate the kLa value of a fermentor. According to Van't Riet (1979), the use of commercial available electrodes with a response time of 2 to 3 seconds should enable a kLa of up to 360 h"1 or 0.1 s 1 to be measured with little loss of accuracy (Stanburg and Whitaker, 1984). The estimation of kLa 35 by this technique depends upon monitoring the increase in dissolved oxygen concentration

of a solution during aeration and agitation. The oxygen concentration of the solution is first

lowered by flushing the liquid with nitrogen gas, until all oxygen is removed from the

solution. The deoxygenated liquid is then aerated until the solution reaches saturation. The

change in dissolved oxygen tension (DOT) is monitored to record its time course.

The actual oxygen concentrations in the broth can be calculated from DOT values.

100%-DOT in water is equivalent to the saturated condition (C*), i.e. the solubility of

oxygen in water. Since air is commonly employed as the source of oxygen, the solubility

of oxygen at 1 atm air in water at 30°C can be determined by using the volumetric

percentage of oxygen in air, i.e. 21% of the value in Table 2.1. The concentration of

dissolved oxygen in the fermentation broth (C/,) can be assumed to be proportional to the

measured DOT value. Integration of equation (2.17) yields:

In {C * -C L) = 0-kLat (2.18)

where 0 is the integration constant. As shown in Figure 2.5, a plot of In (C* - CL) against time (r) yields a straight line with its slope equals to -kta.

2.6.2 Factors Affecting Oxygen Transfer

A number of factors have been demonstrated to affect kLa value achieved in a fermentation vessel. Such factors include the air-fiow rate employed, the degree of agitation, the Theological properties of the culture broth, and the presence of antifoam agents. 36

TABLE 2.1: Solubility of 0 2 at 1 atm pure oxygen in water at various temperatures, and solutions of salt or acid at 25 °C*

Temp, °C Water, O 2 mM/L Temp. °C Water, O2 mM/L 0 2.18 25 1.26 10 1.70 30 1.16 15 1.54 35 1.09 20 1.38 40 1.03 Aqueous solutions at 25 °C Electrolyte 0 2, mM/L cone, M HC1 H2SO4 NaCl 0.0 1.26 1.26 1.26 0.5 1.21 1.21 1.07 1.0 1.16 1.12 0.89 2.0 1.12 1.02 0.71

♦Data from International Critical Tables, vol. Ill, p. 271, McGraw-Hill Book Company, New York, 1928, and F. Todt, Electrochemische Sauerstoffmessuneen , W. de Guy and Co., Berlin, 1958. 37

Slope = - k La

ln(C *-q>

Time

FIGURE 2.5: A plot of the Iti(C*-Cl) against lime of aeration, the slope of which

equals -kj,a. 38 The air-flow rate has a relatively small effect on k,a in conventional agitated systems. The range of air-flow rate employed rarely fall outside the range of 0.5 to 1.5

volumes of air per volume of medium per minute (volumetric air-flow rate, wm).

The degree of agitation has been demonstrated to have a profound effect on the oxygen-transfer efficiency of an agitated fermentor (Peters etai., 1989). The rheology of a

fermentation broth has a marked influence on the relationship between kLa and the degree of agitation. Non-Newtonian fermentation broths present major difficulties in oxygen provision with the result that the productivity of many such fermentations is limited by oxygen availability. Charles (1978) demonstrated that the bacterial cells in a polysaccharide

fermentation made a minimal contribution to the high culture viscosity which was primarily due to the polysaccharide product

Moreover, the high degree of aeration and agitation required in a fermentation frequently gives rise to the undesirable formation of foam. It is necessary to break down a foam before it causes any process difficulties, and this may be achieved by the use of mechanical foam breakers or chemical antifoams. Antifoams were found to decrease the oxygen-transfer rate in solution due to their surfactant nature. For fermentations that are limited by the availability of oxygen, the minimum level of anti foam necessary to control the foaming should be used (Stanburg and Whitaker, 1984).

The volumetric mass transfer coefficient kLa for oxygen in xanthan broth decreases mainly because of the effect of viscosity on intcrfacial area (a). The values of kLa achieved during xanthan gum fermentation in various bioreactors were found to range from 0.1 s 1 at low concentration (< 10 g/L) to as low as 0.01 s 1 at high xanthan concentration (> 20 g/L). Table 2.2 lists some kLa values form various xanthan fermentation and bioreactor studies (Moraine and Rogovin, 1971; Moraine and Rogovin, 1973; Nienow, 1984;

Himmelsbach, 1985; Funahashi e ta i, 1987; Herbst e ta i, 1987; Misra and Barnett, 1987; 39 Pons et a i, 1989; Suh et a i, 1991; Zaidi et al., 1991; Herbst et a i , 1992; Peters et a i,

1992; Tecante and Choplin, 1993; Kessler et a i, 1993; Flores et a i, 1994; Kawase and

Tsujimura, 1994).

2,7 XANTHAN PRODUCTION PROCESS

Currently, commercial production of xanthan gum is from glucose fermentation using the bacterium Xanthomonas campestris. Under growth-limiting conditions, this bacterium produces the gum at the cell surface, secreting it into the surrounding medium, and encapsulating the cell. The gum is produced in large, agitated fermentation units and sent to a holding tank. From there, the gum is subjected to alcohol precipitation (using up two to three volumes of isopropanol to each volume of the fermentation broth), dried, milled, and packaged. The typical industrial xanthan gum production process has already been shown in Figure 1.1 (Margaritis and Pace, 1985).

The conventional batch process for xanthan gum manufacturing is expensive because yield limitation is encountered at high gum concentration. Highly viscous xanthan broth shows good movement in the region of the impeller where shear rates are high.

However, away from the impeller the movement decreases due to the apparent increase in viscosity resulting from the pseudoplastic nature of the fluid (Sittig, 1983). At this point, broth viscosity increases so that the mass transfer, mixing, and heat transfer characteristics of the fermentor deteriorate (Peters et ai, 1989; Gonzales et a i, 1989). High viscosity also decreases the pumping capacity of the impeller which further compounds the problem

(Solomon e ta i, 1981). Therefore, maximum xanthan concentrations achieved in stirred- tank reactors are typically limited to 3.0% (wt/v). High recovery cost is another hurdle to 40 overcome, because the volume of alcohol required for xanthan precipitation is twice the

volume of xanthan broth.

Extensive efforts to improve the production of xanthan gum have been made in the

last two decades. Basic microbial studies include investigations of the biosynthesis

mechanism of xanthan gum (Sutherland, 1977; Ielpi, e ta i , 1981; Jarman and Pace. 1984),

strain selection (Cadmus e ta i, 1976; Whiefield e t a i , 1981), and medium formulation

(Moraine and Rogovin, 1971; Davidson, 1978; Souw and Demain, 1979; Behrens e ta i,

1980; Kennedy and Bradshaw, 1984; Pinches and Pattern, 1986). Environmental factors,

such as pH, temperature, agitation speed, and dissolved oxygen concentration, have been

investigated to further understand the kinetics of substrate consumption and product

formation (Moraine and Rogovin, 1973; Kang and Cottrell, 1979; Weiss and Ollis, 1980;

Peters e ta i, 1989; Shu and Yang, 1990; Vashitz and Sheintuch, 1991; Peters e ta i , 1992).

There have been many attempts to increase xanthan concentration and to lower the energy costs by using new types of bioreactors or agitation designs with improved aeration and oxygen transfer for viscous xanthan fermentation. Emulsion fermentations for xanthan production as described in various patents (Maury, 1982; Engelskirchen etai., 1986;

Viehweg, 1989) have proven to provide better aeration than the conventional process in terms of aeration capacity (Schumpe e ta i, 1991; Ju and Zhao, 1993). The downstream recovery process, however, is complicated by the difficulties in emulsion breaking and surfactant removal (Schumpe et a i, 1991). Other bioreactors, such as stirred tank Rushton turbine fermentors, stirred tank fermentors with draft tubes, bubble column bioreactors, airlift reactors, and fermentors with twin impellers have also been tested for their potential improvements in aerating highly viscous xanthan fermentation broth (Herbst e ta i, 1989;

Suh e t a i , 1991; Nakajima e ta i, 1990). A transport controlled bioreactor using immobilized cell technology for simultaneous production and concentration of xanthan gum 41 was studied by Robinson and Wang (1988). The final concentration of xanthan gum in these works reached an upper limit of -5% because of oxygen transfer limitations in highly

viscous solutions. Several bioreactor concepts that have been developed are discussed in the following section.

2.7.1 Bioreactor Types

It is generally agreed that a mechanically agitated fermentor is required to achieve good heat and mass transfer at high viscosities (Margaritis and Zajic, 1978; Pace and

Righelalo, 1980) and is supported by the data of Bhavaraju et a i (1978) and Henzlcr

(1981). Recent efforts to improve xanthan fermentation mainly focused on the type of impellers used as well as the oxygen transfer performance by each type of the impellers.

However, since the limitations due lo aeration and agitation difficulties were observed in agitated fermentors, other options were also employed to improve xanthan fermentation.

The characteristics and performance of various bioreactor designs and systems reported in the literature regarding xanthan gum fermentation, including stirred-tank bioreactors, foam fermentors, bubble column and air-lift bioreactors, plunging jet bioreactors, and water-in- oil emulsion systems, are summarized in Table 2.2.

As shown in Table 2.2, the performance of stirred-tank bioreactors with various impellers is limited because of the formation of stagnant zone as xanthan concentration exceeds 20 g/L, which subsequently leads to a high power consumption and torque oscillation. Other types of bioreactors, such as foam fermentors, bubble column or air-lift bioreactors, and plunging jet bioreactors, are not applicable in industries because of the difficulties in process scale-up. The water-in-oil emulsions, on the other hand, are able to provide high liquid-liquid interfacial area that consequently benefit the oxygen transfer at TABLE 2.2: Comparison of different bioreactor systems for xanthan gum fermentation.

Max. Xanthan Ferm. OTR kta Cone. Time Productivity ______System Type ______(g/Lh) ( s 1) (g/L) (h) (g/g _Cell/h) ______Remarks Reference

Stirred-Tank Bioreactor

Double disc turbine 0.6 0.03 22-30 80 0.1-0.2 High OTR at low xanthan Moraine & Rogovin, 1971 impeller conc. Moraine & Rogovin, 1973 Low torque fluctuation Funahashi et al., 1987 Low power drop Peters etai., 1992 Stagnant zone in wall region Flores etai., 1994

45° pitch downward- 0.05 20 56 — Better OTR than disc turbine Nienow, 1984 pumping 6-bladed agitator to Stagnant zone still existed with disc turbine 0.1

Intermig impeller 0.7 0.02 50 57 0.107 Reach high xanthan conc. in Himmelsbach, 1985 to short fermentation time Peters et aL, 1992 0.2 High power input Herbst et al, 1992

Intermig impeller with 0.36 0.01 20 135 0.1 Ineffective at low viscosity Herbst et al., 1987 internal draft tube to Advantage at high viscosity 0.1 Reached low xanthan conc.

Twin impeller — — 23-27 50 0.2 Large high shear region Nakajima et al., 1990 High specific production rate High power consumption

Lightnin A-315 impeller — — 35 — — Severe torque oscillation Galindo & Nienow, 1992 at high xanthan conc.

Scaba 6SRGT agitator — — 35 — — Improved bulk mixing Galindo & Nienow, 1993 Instability in torque TABLE 2.2: (corn.)

Max. Xanthan Ferm. OTR kLa Conc. Time Productivity System Type______(g/Lh) (s*) (g/L) (h) (g/g Cell/h) ______Remarks Reference

Helical ribbon screw — 0.01 30 No dead zones throughout the Tecante A Choplin, 1993 impeller length of the impeller Limited capacity to disperse the bubbles

Foam Fermentor — 0.01 20 120 High specific growth rate Misra & Barnett, 1987 to Final yield only 70% 0.1 Need surfactant

Bubble Column — 0.01 20 120 Homogeneous mixing Pons e ta i, 1989 High specific growth rate Suh e ta i, 1991 Hold only small volume Very difficult to scale-up

Air-Lift Bioreactor

External circulation loop — 0.01 25 36 0.05 Low power input Suhefo/., 1991 Poor OTR at high viscosity Kessler et ai, 1993 Low production rate & yield Hard to scale-up

External circulation loop — 0.02 Better OTR than with external Kawase & Tsujimura, 1994 with short draft tubes circulation loop only covered by perforated plates Still needs improvement TABLE 2.2: (cont.)

Max. Xanthan Ferm. OTR Conc. Time Productivity System Type______(g/Lh) (s-*) (g/L) (h) (g/g Cell/h) ______Remarks______Reference

Plunging Jet Bioreactor 0.5 0.02 20 100 --- Very low power input Zaidi etai., 1991 High specific growth rate Poor mixing near wall Hard to scale-up

Water-in-Oil Emulsion

isoparaffm — — 50 0.075-0.1 High liquid-liquid interfacial Schumpe era/., 1991 area Low power requirement Agglomeration at high xanthan conc. Contamination by emulsifier

n-hexadecane — — 65 100 0.13 Fine water-in-oil dispersion Ju & Zhao, 1993 Hard to recover product

Cell Immobilization

Entrapment using porous — — 55 120 Reached high xanthan conc. Robinson & Wang, 1988 support particles Only 25% total xanthan was produced at 55 g/L Oxygen transport limitation Hard to recover product

£ 45 high xanthan concentrations. However, the contamination caused by the use of emulsifier

and the difficulty in product recovery need to be overcome. The immobilized cell

bioreactor can be used for long-term continuous xanthan production if the oxygen limitation

caused by the support material can be eliminated.

Enhanced oxygen transfer and conditions for improved mixing in a centrifugal film

bioreactor, particularly of very viscous liquid media, have been reported (Roubicek and

Feres, 1987; Long and Roubicek, 1988). The use of centrifugal bioreactors, developed for

bioseparation, has been the subject of recent studies for applications in enzyme and

fermentation technology (Parts and Elbing, 1975; Setford e t a i , 1994). Therefore, the

centrifugal bioreactor becomes a potential candidate for the oxygen-limited xanthan

fermentation to reach higher productivity and xanthan concentration than existing systems.

Other bioreactors with improved oxygen transfer efficiency may also be employed for

xanthan gum production, including liquid-impelled loop reactor (Sonsbeek et a i, 1992),

deep jet bioreactor (Moser eta i, 1991), multicompartment bioreactor (Shuler e ta i, 1991),

and downflow liquid jet loop reactor (Li et ai, 1991).

2,7,2 Fermentation Processes

Xanthan gum is usually produced batchwise in stirred tank biorcaetors (Weiss and

Ollis, 1980; Quinlan, 1986; Peters et a i , 1992). As described in previous sections, the conventional batch process for xanthan gum production is expensive. Silman and Rogovin

(1972) studied xanthan gum fermentation in a single-stage continuous process and showed that production of xanthan gum was a function of dilution rate and pH. Continuous production of xanthan gum at steady state conditions was also studied (Vashitz et a i, 1988;

Vashitz and Shcintuch, 1991). The yield and production rate in a continuous process were 46 found to be significantly better than in a batch operation. However, the final concentration of xanthan gum in a continuous operation was lower than that in a batch operation, resulting in a higher product recovery cost

In order to reduce observed strain variations in both batch and continuous fermentations (Cadmus et al., 1976), a repeated-batch process was developed by replacing

80% of fermented broth with fresh pre-sterilized media (Qadeer and Baig, 1987). The efficiency of the culture remained almost unchanged in three cycles of 50 hours each during repeated batch process. The repeated-batch operation was able to not only prevent unexpected contamination but also retain the productivity of the culture. This observation provides a clue that it is possible to reuse the cells of X. campestris without losing their productivities if the cells can be immobilized onto a suitable carrier.

2.7.3 Recovery and Purification

The recovery of xanthan gum from the fermentation hroth includes the following steps: (i) cell removal: heating of broth to lyse cells; (ii) isolation of exopolysaccharide: precipitation of xanthan with different nonsolvents (e.g., alcohols) or multivalent salts; (iii) dewatering and drying: pressing or centrifugation before drying (excessive heat will cause poor color and solubility, and degradation), then drying by use of forced air or vacuum continuous or batch dryers; (iv) milling and packaging; and (v) additional processing.

The critical step in this recovery process with respect to cost is the precipitation of the polysaccharide with alcohol. Energy consumption in this step can be very large, because the alcohol used in xanthan precipitation needs to be recovered by distillation for solvent reuse (Gonzales et a i, 1989). 47 2.7.3.1 Alcohol precipitation

The cost of recovery of microbial polysaccharides, including concentration,

isolation, and purification, is a significant part of the total production cost due to the dilute

nature of the stream leaving the fermentor (about 20 to 30 kg m 3), the presence of

contaminating solids (e.g., cells) and solutes in the stream, and high broth viscosity.

Xanthan is isolated on a commercial scale by precipitation with isopropanol, mainly

because this method complies with the FDA description for the production of food grade

material (Smith and Pace, 1982).

Margaritis and Pace (1985) indicated that the volume of alcohol required for

precipitation of xanthan is relatively independent of polymer concentration but is affected

by the concentration of certain salts. Therefore, it was suggested that increasing the final

concentration of product during fermentation and increasing the salt content of the broth

prior to precipitation will decrease the amount of alcohol used. Also, adjustments in the

fermentation medium and conditions can be used to minimize coprecipitation of salts and

proteins and improve final product purity.

In using alcohol precipitation, the salt type and alcohol concentration are the most

important variables affecting the recovery of xanthan from aqueous solution. The effect of

ethanol concentration on xanthan recovery is expected, because ethanol decreases the

dielectric constant of the solvent mixture, increases the intermolecular attraction between solute molecules (Green and Hughes, 1955), and enhances the binding of cations to the

xanthan molecule (Smith and Pace, 1982). The recovery of xanthan was doubled with an increase in ethanol level from 30% (v/v water) to 65% (v/v water) (Gonzales et al ., 1989).

Increasing salt concentrations up to 1% enhanced the effect of higher ethanol levels but

further increases do not contribute to increased xanthan yields. Potassium chloride was significantly superior to sodium chloride in this respect. 48 According to Gonzales et al. (1989), some advantages can be gained by using

ethanol as compared with isopropanol. For example, ethanol is equally acceptable for food

use and the distillation of ethanol-water systems is better characterized than any other

alcohol-water system. Average xanthan recoveries of 88-90% were achieved under optimal conditions (65% ethanol with 1% KC1 and a temperature of 15°C). However, the cost of ethanol is similar to that of isopropanol (about $1.25 per gallon), and this represents a

major part of the production cost.

2.7.3.2 Freezing-thawing method

The freezing-thawing method was described in a European patent (Haze et a !.,

1989) and was claimed that a large amount of water contained in a hydrated solution of a hydrophilic macromolecular substance can be easily removed by freezing it and then immersing and thawing it in a hydrophilic organic solvent.

Xanthan solution was first placed in a freezer at -20°C to obtain a frozen product.

This frozen product was then immersed and thawed in a 2-fold amount (v/w) of ethanol at room temperature. Hydrophilic organic solvents, such as methanol, ethanol, isopropyl alcohol, and acetone, which have water solubility exceeding 25% and a lower boiling point than water were suggested for actual use. The product was dehydrated using a centrifuge and vacuum-dried at 60°C to yield a dry product with a porous structure. A dehydrated xanthan gum with solid content exceeding 25% was obtained in 6 hours. The used organic solvent may be recovered by distillation and the recovered solvent is expected to be recycled to use in a subsequent process. However, this dehydration process has the same problem as the use of ethanol lo recover xanthan gum - the vast costs in both the amount of ethanol used and the facilities for distillation. 49

2.7.3.3 Other purification methods:

Multivalent metal salts have been used as an alternative to alcohol precipitation hut

the xanthan product obtained was found to be insoluble, and an appreciable amount of the

salt must be removed before the product can be considered suitable for food use (Kennedy

and Bradshaw, 1984).

Reported methods used in the isolation of xanthan from fermentation broth

(Gonzales etai., 1989) include: ethanol (Funahashi e ta i, 1987; Lehmann, 1985), ethanol with KC1 (Stauffer and Leeder, 1978; Thonart e ta i , 1985), ethanol with \ % KC1 (Cadmus e t a i , 1978), ethanol (multistep) with 1% KC1 (Jeanes e ta i, 1961; Jeanes et a i, 1976;

Souw and Demain, 1979; DeVuyst e ta i, 1987), ethanol (multistep) with NaCl and EDTA

(Holzwarth, 1978), methanol (Moraine and Rogovin, 1973), methanol with 2% KC1

(Moraine and Rogovin, 1971; Rogovin e ta i, 1961), isopropanol (Kennedy et a i, 1982;

Jarman and Pace, 1984, Thome et a i, 1988), isopropanol with salts (Garcfa-Ochoa et ai,

1993), acetone (Davidson, 1978), acetone with saline (Tait et a i, 1986) and acetone with

1% KC1 (Pinches and Pallent, 1986), quartolan (Fiedler and Behrens, 1984), and tert- butanol (Flahive e ta i, 1994). However, industrially the large amount of alcohol used in xanthan recovery still represents a major part of production cost. 50

2.8 CELL IMMOBILIZATION

2.8.1 Advantages and Limitations

2.8.1.1 Advantages

From a process engineering point of view, the benefits of immobilized cell systems are high reactor productivity at high throughput and ease of product recovery in the system

(Atkinson, 1986). Black (1986) indicated that the advantages of immobilized cell technology can be viewed as falling into three categories. The first relates to the handling characteristics of the immobilized aggregates, where solid/liquid separation is required.

The second relates to the insensitivity to cell washout in continuous systems. Finally, novel process organization becomes possible, such as the countercurrent flow of biocatalysts (e.g. cells, organelles, and enzymes) and liquid.

These benefits are mainly because of the fact that the microenvironments offered by the carrier maybe more favorable to the organism. For instance, the rate of growth for living cells may be increased by immobilization. Also, because of the possibility of higher cell loading, reaction rates are expected to be higher. Cell immobilization also favors cell separation, enables higher dilution, leads to possible gas evolution in the center of microbial aggregates, allows continuous fermentors to be operated on a drain and fill basis, and greatly facilitates recycling or reuse of microorganisms.

There are other features which have to be considered for the application of cell immobilization, including: the formation of heterogeneous populations of organisms within microbial aggregates; the effects of immobilization on overall rates of growth and stoichiometry; protection against contamination; manipulation of growth rate in continuous systems; independent of dilution rate; manipulation of the cells as a discrete phase; possible 51 use of optimum aggregate sizes leading to maximum microbial activities; and possible

spatial location within reactors of different microbial populations (Atkinson , 1986).

2.8.1.2 Limitations

Nevertheless, cell immobilization systems do suffer from a number of problems.

The exploitation of immobilized cells within a process places a heavy requirement on

maintaining the desired local conditions throughout the process, particularly with regard to

the physical environment and the fluid mechanics, e.g. shear. The initial expense for such

a system is usually high, and the mechanical properties of the system are more complex

than those of free cells. All these criteria have to be taken into account in order to provide a continuous, recycling process and effective agitation (Phillips and Poon, 1988). The microbial aggregates which result from immobilization are often subject to diffusion limitations on the rates of reaction, especially when high molecular weight substrates or products are encountered (Atkinson, 1986).

There are several major problems associated with conventional immobilized cell bioreactors that have prevent their wide industrial applications. For a non-growing, immobilized cell system, the loss of cell viability results in decreasing reactor productivity over time, thus severely limit the operating life of the bioreactor. For an actively growing system, complications in maintaining bioreactor stability make continuous operation of the bioreactor rather difficult at an industrial scale. For examples, packed-bed and membrane bioreactors tend to get clogged by cell biomass and fluidized-bed biorcactors are subject to unstable bed expansion due to biofilm growth. Conventional packed-bed and membrane bioreactors also suffer from high pressure drop and gas entrapment inside the reactor bed that reduces the reactor working volume substantially. Moreover, conventional packed-bed and membrane bioreactors tend to accumulate dead cells over time and thus gradually lose their production capability (Lewis and Yang, 1992). However, judicious selection of the 52 diffusion conditions is important and can lead to enhanced reaction rates which are affected by the nature of the reactions taking place.

2.8.1.3 Changes in cell properties

According to Phillips and Poon (1988), the growth of microbial cells immobilized on a solid earner essentially follows the pattern of free cells. Growth of immobilized cells can be viewed as an alternative to growth of free cells but with a precise control over the process. Cells in the dispersed state are indistinguishable from those in the liquid phase and have low sedimentation velocities requiring centrifugation for cell separation. Such cells maintain high conversion rates, whereas growth in aggregates would result in diffusion limitations (Black, 1986). As described by Black (1986), cell immobilization allows biological particles to be produced incorporating a wide variety of microorganisms.

Accumulation of biomass occurs during the start-up of an immobilized cell system, when the fluid mechanical conditions are changed, or when the prevailing conditions are insufficient to remove excess growth. However, it becomes possible to maintain a constant amount (hold-up) of biomass within a reactor.

Cell concentration in an immobilized cell reactor can be severalfold higher than those in free cell systems. This may more than compensate for any reduction in cell activity caused by diffusional limitations (Black, 1986). The ability to operate without cell growth is an important characteristic of immobilized cell reactors and allows the possibility for continuous production of secondary metabolites with more efficient bioconversions. It has been demonstrated that by choosing a glucose-based medium that will support product formation but not cell growth, the fermentation reaction may be diverted to produce desirable products instead of simply increasing biomass (Foerberg et al., 1983; Mosbach et al., 1983). 53 In an unperturbed environment, the exponential growth phase will finally slow down and reach a stationary phase in which the microbial population remains constant.

Kuu and Polack (1983) indicated that immobilized cells tend to retain their shape and rigidity, although under conditions of rapid growth the size of immobilized cells is generally smaller than that of free cells. Changes in cell metabolism after immobilization were categorized by Phillips and Poon (1988) as follows: 1. the total number and concentration of cells in the matrix increases; 2. cell generation time is usually shorter; 3. changes occur in macromolecular cell components; 4. the pH optimum may shift; 5. the temperature optimum may shift; 6. operational stability usually increases after successful immobilization; 7. the lag time may disappear; and 8. the respiration rate, oxygen consumption, and enzyme activity usually increase.

Direct investigation of the process of immobilization is possible by microscopic observation, especially using scanning electron microscopy (SEM). The image obtained by

SEM is a picture of the surface of the sample with a very great depth of field so that an object can be viewed in perspective. Internal structure of a sample can be observed by

SEM at the fracture surface. Observations under the SEM are used often in the identification of microorganisms and enzymes (Costerton and Marks, 1977) to provide evidence of cell viability (Wada et al ., 1980) and direct evidence of immobilization

(Brooker and Fuller, 1976), and to show the morphology of the carrier (Buchholz, 1979).

2.8.2 Methods of Immobilization

A general definition of immobilization is given by Atkinson (1986) as the presence of organisms, either on surfaces or within particles by means of adhesion, flocculation, or artificial methods, such as covalent bonding. The immobilization of cells, organelles, and 54 enzymes has become a powerful and well-accepted technology with various biotechnological and biomedical applications (DiCosmo et a i , 1994).

It’s a well-known fact that the living microorganism can attach to each other and to solid surfaces in the form of films (Fletcher and Floodgate, 1976; Brooker and Fuller,

1976; Bums, 1979; Knapp and Howell, 1980). The most significant recent advance in immobilization technology is the development of the ability to immobilize any organism independent of its nature flocculent or adherent properties (Atkinson et a i , 1984).

A range of methods for cell immobilization arc now available based upon solid supports and gels. There are three major techniques commonly used for cell immobilization: adsorption, entrapment, and coupling. The selection of the immobilization technique depends on the applications. Figure 2.6 compares the advantages and disadvantages of these cell immobilization techniques.

2.8.2.1 Adsorption (Kolot. 1988)

A great variety of different microorganisms such as species of Bacillus and

Pseudomonas can adsorb to wood, glass, ceramic, and plastic. Some microbes can attach because of the formation of an adhesive disk to anchor the cell to the support. For others, a charge on the cell surface and cell wall composition provide the necessary eloctrosialic and ionic sites for attachment to the carrier. Immobilization by adsorption has several advantages compared with other techniques: the cells remain alive and their enzymatic activity is not affected. The only disadvantage of this method is that electrostatic interactions are influenced by pH changes occurring during fermentation. IMMOBILIZATION 1. Facilitate separation 4. Increase reaction rate 2. Prevent cell washout 5. Increase permeability 3. Increase cell density 6. Cells can multiply

Adsorption Entrapment Coupling (Electrostatic interaction, (Partial covalent bond) (Covalent bond)

A d v a n ta g e : A d v a n ta g e : A d v a n ta g e : 1. Cells remian alive 1. pH is not altered 1. No diffusion limit 2. Enzymatic activity 2. No +/- accumulate Disadvantage: Disadvantage: Disadvantage: 1. Chemical change 1. Influenced by pH 1. Diffusion problem 2. Less enz. activity

FIGURE 2.6: Summary of cell immobilization techniques.

LA La 56 2.5.2.2 Entrapment (Kierstan and Coughlan. 1985)

Polyacrylamide, the most commonly used matrix for the entrapment of biocatalysts,

has the property of being non-ionic. Thus, the pH profile characteristics of the free

cells/enzymes are minimally altered. Moreover, charged substrates and products do not

accumulate nor are they depleted in the matrix. However, the failure of the matrix to

interact with the entrapped cells/enzymes does little to prevent leakage. Therefore,

generation of a high degree of cross-linking is necessary to obviate this problem.

Unfortunately, there are diffusional problems arising from the use of highly cross-linked

matrices especially where large substrate are concerned.

2.8.2.3 Coupling (Kolot. 1988)

The mechanism of covalent cross-linking is based on covalent bond formation

between activated inorganic support and cells, and requires the use of a binding agent. To

introduce the covalent linkage, chemical modification of the carrier surface is necessary.

The advantage of this system compared with entrapment is that it is free from the diffusion

limitation. Unfortunately, coupling agents usually are toxic and cells can retain only 60-

75% of their enzymatic activity.

2.8.3 Immobilized Cell Bioreactors

Immobilized cell reactors have been used successfully for a long time in wastewater treatment and in the vinegar industry. Modern developments of immobilized cell technology started in Japan in the 1970's with processes such as the production of L- aspartic and L-malic acids by immobilized cell processes (da Fonseca et a l, 1986). The 57 concept of cell immobilization is intrinsically linked to continuous operation even though the majority of all current bioreactors are batch stirred tanks.

Free cell systems can be modified to exhibit some of the characteristics of immobilized cell systems. Such characteristics may be achieved using cell retention by ultrafiltration, especially with hollow fiber membranes (Kan and Shuler, 1978; Vickroy et al., 1982; Inloes et al, 1983) and other membranes (Margaritis and Wilke, 1978). Cell settlement followed by recycle is another method of achieving high cell densities in fermentors. This concept has been used in wastewater treatment (da Fonseca et al., 1986),

The development of cell recycle systems for the continuous production of ethanol using yeast have also been achieved (Cysewski and Wilke, 1977; Fricker, 1983; Comberbach and Bu'lock, 1984).

In 1983, the Japanese Research Association for Petroleum Alternative Development

(RAPAD) developed two continuous pilot plant immobilized whole cell processes for the production of fuel ethanol from molasses (Oda et al., 1983). Such a novel fixed bed reactor, in spite of its reduced area when compared to particulate beds, can give high productivities. The improved performance is mainly due to its open structure which allows good gas release and prevents blockage (da Fonseca et ai, 1986).

Another design is the rotating disc fermentor, a film fermentor developed for the cultivation of mycelial fungi (Blain et al., 1979). The mycelia grow on the surfaces of a series of vertical polypropylene discs which rotate half submerged in the culture medium.

The system can be operated aseptically and has been tested using a wide range of filamentous fungi. By encouraging the passive immobilization of the fungi, disc fermentors offer potential for the exploitation of filamentous organisms which are associated with high non-Newtonian viscosities in suspended cultures. Other advantageous features include low power input requirements, ease of biomass removal. 58 and relative simplicity of construction (da Fonseca et al, 1986). Rotation film fermentors have shown good performance in the production of citric acid by A. niger (Anderson et al,

1980).

Continuous fermentations in spiral wound fibrous bed bioreactors were successfully developed by Lewis and Yang (1992) in the production of propionic acid from lactate, and more recently from cheese whey permeate (Yang et al ., 1994). These reactors have also been employed with yeast (Shu, 1992) and animal cells for recombinant protein production (Zhu, 1995). The use of fibrous matrices as carriers for cell immobilization by natural attachment is viewed as an effective method to rapidly reach high cell density and to improve reactor productivity of a fermentation process. It facilitates the separation of cells from products in solution and allows reactor operation at high dilution rates without cell washout. Also, the fibrous bed reactor was able to operate for a long term without encountering clogging, degeneration of cells, or contamination problems. This type of immobilized cell bioreactor opens a new approach towards the production of other industrial fermentation products.

Xanthan gum fermentation using cells immobilized in porous beads to control broth viscosity in the fermentor was studied by Robinson and Wang (1988), The entrapment technique was used by adding a liquid culture to dry autoclaved porous support particles

(celite beads). A final xanthan concentration of 5.5% in the particles was obtained.

However, xanthan concentration in the bulk fermentation broth was only -3% and only

25% of total xanthan was produced at the maximal internal concentration with a significantly lowered productivity. It was found that the total xanthan accumulation in the system was very sensitive to oxygen transport limitation. Therefore, it is noted that to reach a high cell density without encountering severe oxygen limitation to the immobilized 59 cells will be a major concern in choosing a proper immobilization technique for

Xanthomonas campestris.

2.9 ULTRAFILTRATION PROCESSES

Fermentation broths are complex mixtures of biomass, dissolved macromoleculcs, and electrolytes (Fane and Radovich, 1990). Separation and concentration of these materials is difficult because the desired products are usually in dilute solution, with similar physical and chemical properties.

Membrane processes, which serve as a molecular sieve to separate solute molecules of different molecular size, provide a means of separation and concentration at the molecular and fine-particle level. Industrial applications of various membrane separation technologies have been used in various fields, for example, protein purification and waste water treatment (Shuler and Kargi, 1992; Kimura, 1992). Depending on molecular size cut-off, different membranes can be used for the separation of different molecules.

2.9.1 Significance and Applications

Table 2.3 shows the characteristics of various membrane processes. To select a suitable process from these membrane separation technologies, it is very important to distinguish the usage between them.

According to Cheryan (1986), reverse osmosis retains all components other than the solvent. Ultrafiltration, on the other hand, is designed to retain only macromolecules or particles larger than about 10-200 A (about 0.001-0.02 Jim), which covers molecules that 60

TABLE 2.3: Characteristics of some membrane processes. | Process Driving Force Permeate Retentate*

Osmosis Chemical Potential Water Solutes

Dialysis Concentration Water + Large Molecules

Difference Small Molecules

Electrodialysis Electromagnetic Water + Non ionic Solutes

Field Ionic Solutes

Reverse Osmosis Pressure Water Solutes

Ultrafiltration Pressure Water + Large Molecules

Small Molecules

Micro filtration Pressure Water + Large Suspended

Dissolved Solutes Particles ♦Also includes water. (Adapted from Cheryan, 1986) 61 range from about 1000 in molecular weight to about 1,000,000. The membranes used for ultrafikration are finely microporous and asymmetric. Microfiltration is an extension of ultrafiltration, but the membranes have a larger pore size. Thus microfiltration retains suspended particles in the range of 0.10 ^im to about 10 Jim. Membranes that arc very finely microporous (less microporous than ultrafiltration membranes) are used in dialysis

(Fane and Radovich, 1990). The operation mechanism that mainly distinguishes the more common membrane processes - microfiltration, ultrafiltration, and reverse osmosis - is the application of hydraulic pressure to speed up the transport processes.

Some typical examples of components that fall under these processes are given in

Figure 2.7. Because of the selectivity of the membrane, specifying which component permeates and which component is retained, ultrafiltration can be viewed as a method for simultaneously purifying, concentrating, and fractionating macromolccules or fine colloidal suspensions.

Ultrafiltration is operated by pumping the feed solution under pressure over the surface of a suitably supported membrane with appropriate chemical nature and optimum physical configuration. In this process, the pressure gradient across the membrane would force solvent and smaller species through the pores of the membrane, while the larger molecules would be retained. The "retentate" stream will thus be enriched in the retained macromolecules, while the permeate stream will be depleted of the macromoleculcs.

Because ultrafiltration deals with the separation of fairly large molecules, the osmotic pressures involved in the processes are relatively low. Thus ultrafiltration offers the advantage of needing low pressures for operation, which would lower equipment and operating costs by a considerable margin. A further advantage of ultrafiltration as compared to conventional dewatering processes, such as evaporation and freeze concentration or freeze drying, is the absence of a change in phase or state of the solvent 62

MOLECULAR SIZE EXAMPLE MEMBRANE PROCESS WEIGHT

-100 pm POLLEN-

-10 pm STARCH —

BLOOD CELLS— TYPICAL BACTERIA— MICRORLTRATION —1 pm

SMALLEST BACTERIA—

-1000 A

DNA, VIRUSES —

100,000 — -100 A ALBUMIN —

10,000 — ULTRAFILTRATION VITAMIN B12 — 1000 — - 1 0 A r t 1 f f C C

WATER — REVERSE OSMOSIS - 1 A NaCI —

(Adapted from Cheryan, 1986)

FIGURE 2.7: Examples of components separated by MF, UF, and RO processes. 63 during the process, and this should result in modest energy demand. The schematic

diagram and the solute concentration profile of membrane ultrafiltration process is

illustrated in Figure 2.8.

However, there are some limitations to ultrafiltration process, including the low

mass transfer rates obtained with concentrated macromolecules, and the high viscosity that

makes pumping of the retentate difficult. A major limiting step in the use of pressure-

driven membrane processes is the "fouling" of the membrane. Fouling manifests itself as a

decline in flux with time of operation, and the flux decline occurs when all operating

parameters, such as pressure, flow rate, temperature, and feed concentration, are kept

constant. Detailed reviews related to fouling problems is discussed later in Section 2.9.3.

2.9.2 Theory

Description of performance in axial flow hollow fiber modules has been reported widely, especially in the application of protein ultrafiltration. However, in the area of polysaccharide ultrafiltration no such study has been published yet. The filtration performance is usually analyzed in terms of volumetric flux of permeate and pressure drop.

Volumetric flux is the membrane capacity per unit area and unit time, and is determined by measuring the permeate flow rate and dividing it by the cartridge membrane area.

It is a well-known fact that the flux of a solution is substantially lower than the flux of pure water. This has been explained by membrane scientists in several ways; by the gel layer model, the osmotic pressure model, and the resistancc-in-series model (Yeh and

Cheng, 1993). The application of the different models depends primarily on the diffusivily

(i.e. the size) of the solutes. In contrast to small molecules, such as salt and sugar, the diffusivity of large solute molecules, polymeric and colloidal substances, for example, is UF membrane

(a)

Gym : limiting (gel) concentration

Cb : bulk solute concentration lim Cp ; permeate solute concentration

J : solvent flux

J s : solute flux

(b)

FIGURE 2.8: Schematic diagram and concentration profile of UF process. 65 very limited. This means that once a large molecule has been transported to the membrane surface the chances of it returning to the bulk solution by diffusion are substantially reduced (Jftnsson, 1993). The concept of concentration polarization , referring to the development of concentration gradients close to the membrane surface caused by the different rates of transport of various components, can be introduced here to explain the formation of film resistance. For dialysis, concentration polarization decreases the effective concentration driving force, whereas for reverse osmosis, ultrafiltration, and microfiltration, the effect is to increase the concentration presented onto the membrane surface. Thus, the film resistance to mass transfer in the liquid at the membrane surface may be significantly increased, and the flux may be significantly decreased (Fane and

Radovich, 1990).

Of the above mentioned transport models, the resistance-in-series model is the most generally applicable one (Yeh and Cheng, 1993). The resistance-in-series model is based on the assumption that there arc several distinct resistance in series which control the volumetric flux, Jv (in m3-m-2-s_1).

A P~ An J - (2.19)

in which A P and An are transmembrane pressure difference and osmotic pressure difference across the membrane (in Pa), t] is the solution viscosity, and ^ R , is the resistance to transport and is mainly the sum of Rm, intrinsic resistance of membrane (Pa- nr3 m2 s), and /?/, film resistance owing to concentration polarization, solute adsorption, pore blocking, and membrane fouling (Pa- m 3 m2 s).

AP can be determined by

(2.20) where Pi and P0 stand for the inlet and outlet pressure of the tubeside (in Pa), respectively,

and Pp is the permeate pressure of the shellside (in Pa). The osmotic pressure difference. A n , is determined by the nature of the solutes in solution and their concentration

differences across the membrane.

A n = n (Cm) - n (Cp) = 7r(C J (2.21)

where 7t(Cm) is the osmotic pressure of solution at the inside membrane surface and ftC p)

is the osmotic pressure of permeate solute concentration. According to the van’t Hoff

equation, the relationship between the osmotic pressure k and the solution concentration C,

(in g/L) is:

(2.22) where R is the gas constant, 8.314 J/Kmol, T is absolute temperature, °K, and M is molecular weight. It can be seen that the osmotic pressure is proportional to the concentration and inversely proportional to the molecular weight (Mulder, 1991), In the case of large molecule solutions, the value of i^Cj) is relatively small compared to the transmembrane pressure.

According to equation (2.19), the effective pressure is reduced as the osmotic pressure of the retentate increases, and will lead to decreased permeate flux. When ultrafiltration is carried out for pure water instead of a solution, there is no osmotic pressure present. Additionally, for a fresh membrane before solute adsorption and fouling, the volumetric flux is related with Rm, and equation (2.19) then reduces to

(2.23) 67 Therefore, the value of Rf can be determined from equation (2.19) upon knowing the values of other parameters computed from equations (2.20) through (2.23) as

, AP - Att AP , Jv = ------= ------when AP » An and R f » Rm (2.24) r}(Rm + Rf ) rj Rf

Since almost all feed components will foul the membranes to a certain extent, some understanding is required to control the fouling process and to improve the process efficiency.

2.9.3 Factors Affecting Ultrafiltration

Fouling is a result of concentration polarization, specific interactions between the membrane, and various solutes in the feed stream and between adsorbed solutes and other solutes in the feed stream. It is difficult to establish general rules or theories about the nature and extent of fouling that will be universally applicable. However, various factors have been found to have significant effects on the performance of ultrafiltration.

2.9.3.1 Physico-chemical factors

Fane (1986) has discussed initial ultrafiltration flux decline in terms of the boundary layer theory, adsorption, and pore plugging. Long-term flux decline was attributed to deposition/adsorption and cake consolidation. Each component of a feed stream will react differently with the membrane, and the conformation, charge, zeta potential, and other factors will have a significant bearing on these membrane-solute interactions. Proteins, lipids, salts, and microorganisms in the feed stream have been frequently mentioned in the literature as affecting membrane flux. 68

2.9.3.2 Feed concentration

In addition to the complicated interactions of feed components, process parameters

such as temperature, flow rate, pressure, and feed concentration, as well as overall

equipment design, have great influence on concentration polarization. As the concentration

of solids in the feed increases, its viscosity and density increase and its diffusivity

decreases. These changes in the physical properties will affect the absolute value of the

flux, but higher feed concentrations will also usually aggravate film resistance.

2.9.3.3 Shear rate

The feed velocity or shear stress at the membrane surface is a very important factor

influencing membrane flux. High shear rates generated at the membrane surface tend to

shear off deposited material and thus reduce the hydraulic resistance of the fouling layer.

Severe decreases in flux can be observed at too low velocities. The most common method

of generating high shear rates or turbulence needed to minimize the thickness of the gel

layer is to increase fluid velocity or the recirculation rate and/or decrease flow channel

dimensions.

2.9.3.4 Viscosity

The volumetric flux generally decreases rapidly with increasing the solution

viscosity, mainly because of the increased film resistance caused by concentration

polarization (Wang, 1988; Gill e ta i, 1988). However, the xanthan solution shows a high

degree of pseudoplasticity, i.e., the viscosity decreases rapidly as the shear rate increases

(Kang and Pettitt, 1993). This shear-thinning property should allow efficient ultrafiltration of xanthan polymer at high shear rates. 69

2.9.3.5 Transmembrane pressure

It has been reported by Cheryan (1986) that when transmembrane pressure is in the region where no gel is formed on the membrane, the flux increases as pressure increases, though usually not linearly for macromolecular feeds. As the pressure increases further, the concentration polarization layer reaches a limiting concentration and the flux becomes independent of pressure and becomes mass-transfer limited. Any pressure increase beyond this point will only increase the flux momentarily, but as soon as the equilibrium is reestablished between the rate of transport of solute to and from the membrane surface, the flux remains essentially unchanged. Increasing pressure above a critical point may result in a lower flux, because at high pressures the gel layer gets compacted and becomes less permeable to water. Bruin et al. (1980) also observed that the specific resistance of the gel increased with decreasing shear at the wall and with increasing pressure in skim milk processing.

2.9.3.6 Temperature

The effect of temperature on the performance of ultrafiltration is not too clear.

Increasing temperature is expected to result in higher flux by lowering viscosity and increasing diffusivity. However, this is generally harmful for biological systems, and heat damage, such as protein denaturation, will result in a lowering of the flux (Kuo and

Cheryan, 1983). A high temperature, however, may not be a problem to xanthan gum since it is thermally stable for up to ~80°C.

2.9.3.7 Other factors

Electric fields have been used to counteract polarization in conventional filtration

(Bier, 1971), and more recently combined with ultrafiltration membranes to reduce fouling 70 (Henry Jr. et a l, 1977; Chowdia et a l, 1981; Hong and Lee, 1986; Wakeman and

Tarleton, 1987; Bowen and Sabuni, 1992) and thus to allow operation at higher permeation rates. In these applications the electric field is arranged so that the retained solute tends to electrophorese away from the membrane. This either prevents formation of a dense gel layer or makes the gel layer more permeable. The electro-ultrafiltration modules can then be designed to concentrate or to fractionate polyelectrolytes. If the electric field also drives an electroosmotic flow across the membrane, this can increase permeation rates severalfold.

2.9.4 Membrane Selection

The following characteristics are important in determining the suitability of a membrane for separation applications: (1) porosity, (2) morphology, (3) surface properties,

(4) mechanical strength, and (5) chemical resistance. Properties such as resistance to compaction under pressure, cleaning chemicals, bacterial degradation, and temperature arc also important for industrial use. In actual practice the achieved separation is a function of intrinsic membrane properties, process operating conditions, and module geometry

(Kulkami e ta l , 1992).

Table 2.4 lists various polymeric and inorganic materials for ultrafiltration membrane manufacture (Kulkarni et al., 1992). Most ultrafiltration membranes are polymeric in nature, although recently inorganic membranes have also become available.

The formation of an asymmetric membrane structure, i.e., an upper skin that is permselective and a more porous substructure for mechanical support, is an important element for ultrafiltration membranes. While many polymers have been examined for use as membrane materials, only a few are widely used (Lloyd and Meluch, 1985). The most common ultrafiltration membrane is based on polysulfone, characterized by having in its 71

TABLE 2.4: Typical Ultrafiltration Membrane Materials.

Polymeric Inorganic Polysulfone y- Alumina/a- Alumina Polyethersulfone Borosilicate glass Cellulose acetate Pyrolyzed carbon Regenerated cellulose Zirconia/Stainless steel Polyamides or Zirconia/Carbon Polyvinylidenefiuoride Polyacrylonitrile

(Adapted from Kulkami et al ., 1992)

FIGURE 2.9: Structure of polysulfone (Adapted from Leslie et al., 1974). 72 structure diphenylene sulfone repeating units. The structure of polysulfone is given in

Figure 2.9 (Leslie e ta i, 1974).

Polysulfone membranes are useful for ultrafiltration applications due principally to the following characteristics (Cheryan, 1986): (a) wide temperature tolerance, which typically holds up to 75°C; (b) wide pH tolerances: polysulfone can be continuously exposed to pH values from 1 to 13, an advantage for cleaning purposes; (c) fairly good chlorine resistance usually permitting the use of up to 200 ppm chlorine for short-term sanitation, and up to 50 ppm chlorine for long-term storage of the membrane; (d) easy to fabricate in a wide variety of configurations; and (e) wide range of pore sizes available for ultrafiltration applications, ranging from 10 A to 200 A, or molecular weight cut-offs from

1000 up to 500,000 in commercial-sizc modules.

The main limitations of polysulfone membranes are the apparent low pressure limits, typically 25-30 psig with polysulfone hollow fibers. However, industrially applied spiral wound membranes can withstand pressures up to 70-80 psig.

The choice of module is somewhat important to the performance of ultrafiltration

(Fane and Radovich, 1990). Usually it is a compromise between capital cost, operating cost, availability, and performance of a particular membrane. The membrane module is designed to satisfy several criteria, including physical support, packing density, effective fluid management, suspended solids capability, in situ cleaning, and ease of maintenance and replacement. The characteristics of several module concepts that have been developed are summarized in Table 2.5. TABLE 2.5: Membrane module concepts and their characteristics.

Module concept

Characteristic Flat plate Spiral wound Shell and tube Hollow fiber

Packing density Moderate Moderate Low High

(m2An3) (200-400) (300-900) (150-300) (9000-30,000)

Fluid management Good Good High pumping costs Good (with spacers) (with spacers)

Suspended-solids Moderate Poor Good Poor capability

Cleaning Sometimes difficult Sometimes difficult Easy Backflushing possible

Replacement Sheet (or cartridge) Cartridge Tubes Cartridge

{Adapted from Fane and Radovich, 1990) CHAPTER III

IMMOBILIZED CELL XANTHAN GUM FERMENTATION

3,1 SUMMARY

Xanthan gum fermentation using immobilized cells of Xanthomonas campestris in a centrifugal, packed-bed bioreactor (CPBR) was studied. Cells were immobilized in a fibrous matrix by natural attachment to fiber surfaces to increase cell concentration and xanthan production in the bioreactor. The bioreactor was operated as a repeated batch reactor to study the feasibility of long-term xanthan production using the immobilized cells.

The results showed that cells began to be immobilized in the fibrous matrix at 24 hours when xanthan gum was produced and no suspended cells were found in the fermentation broth after 60 hours of cultivation. A cell-free xanthan broth with -85% product yield was obtained. Cells in the bioreactor were repeatedly used for xanthan production and the production rate, as well as xanthan quality, remained almost unchanged in every subsequent batch cycle for the entire period of over 3 weeks studied. The volumetric xanthan productivity in CPBR was -1 g/L h, which was 2-3 times higher than that from conventional batch fermentation in stirred tank reactor (STR) with free cells. The higher productivity was attributed to the higher cell density, -7 g/L, attained in the CPBR.

However, the specific xanthan productivity was lower in CPBR than in STR.

74 75

This was because the relatively low cell viability in CPBR, which might have been caused

by oxygen limitation.

Keywords: xanthan gum fermentation, cell immobilization, centrifugal bioreactor.

3.2 INTRODUCTION

Xanthan gum is a microbial polysaccharide commercially produced by fermentation

with Xanthomonas campestris (Cottrell et al., 1980; Jeans et al., 1961). The present

industrial process for xanthan gum production is energy-intensive and costly. This is

mainly because the high viscosity of xanthan gum causes the agitation and aeration in the

conventional stirred tank fermentor to be extremely difficult and consequently limits the

final xanthan concentration from fermentation to below -3% (wt/v) and productivity to

-0.5 g/L h. There have been many attempts to increase xanthan productivity and to lower

the energy costs by using new agitation designs (Funahashi et a l, 1987; Galindo and

Nienow, 1992; Galindo and Nienow, 1993; Herbst et a l, 1987; Herbst et a l, 1992;

Himmelabach, 1985; Nakajima e ta l ., 1990; Nienow, 1984; Peters et a l, 1992; Tecante

and Choplin, 1993), new types of bioreactors (Kawase and Tsujimura, 1994; Kessler et a l, 1993; Misra and Barnett, 1987; Pons e ta l, 1989; Suh e ta l, 1991; Zaidi e ta l, 1991), or new systems with improved aeration and oxygen transfer (Ju and Zhao, 1993; Robinson and Wang, 1988; Schumpe et a l, 1991). However, none of these fermentation

improvements was able to economically produce xanthan gum for industrial use. A new, innovative production process that is energy efficient and cost effective is thus needed.

Immobilized cell reactors have been successfully used for many years in wastewater treatment and in the vinegar industry. In general, immobilized cell bioreactors have high 76 reactor productivity because of the increased cell concentration (Atkinson etal., 1984;

Black, 1986; DiCosmo etal., 1994). They allow continuous fermentors to be operated on a drain and fill basis, that subsequently facilitates recycling or reuse of microorganisms

(Atkinson, 1986). Xanthan gum fermentation using cells immobilized in porous celile beads to control broth viscosity in the fermentor was studied by Robinson and Wang

(1988). In spite of the limited oxygen transfer rate inside carrier particles, simultaneous production and concentration of xanthan gum, reaching ~5% xanthan in the particles, was achieved. However, separating and recovering xanthan gum from the particles was very difficult, if not impossible.

Enhanced oxygen transfer and conditions for improved mixing in a centrifugal film bioreactor, particularly of very viscous liquid media, have been described by several authors (Long and Roubicek, 1988; Roubicek and Feres, 1987). The use of centrifugal bioreactors, developed for bioseparation, has been the subject of recent studies for applications in enzyme and fermentation technology (Parts and Elbing, 1975; Setford eta!.,

1994). In this work, a centrifugal, fibrous bed bioreactor was developed for viscous xanthan gum fermentation. Fibrous bed bioreactor has been successfully used in fermentation and cell culture, with greatly enhanced cell density, productivity, and long­ term stability (Yang etal., 1994; Zhu, 1995).

During batch fermentation, cells are under a physiological change from trophophasc to idiophase. After entering the idiophase, cells stop growing, but xanthan gum continues to accumulate (Behrens et al., 1980; Pinches and Pattern, 1986). Thus, because of the secondary metabolite nature of xanthan, theoretically as long as the cells were able to be maintained viable at certain levels, the production of xanthan could be expected to last.

Even with reduced viability, the high cell numbers present in the fibrous matrix can compensate the problem. Moreover, fermentation conditions can be shifted from growth- 77 oriented for reactor startup to production-oriented for the following production period. The productivity of xanthan polymers can thus be optimized in the immobilized cell bioreactor.

The high viscosity of xanthan solution at low concentration presents a major challenge in mixing xanthan broth during fermentation (Zaidi eta l ., 1991). The agitation and aeration (oxygen transfer) in a conventional stirred-tank bioreactor are not appropriate for the immobilized cell bioreactor. However, the xanthan solution shows a high degree of pseudoplasticity, i.e., the viscosity decreases rapidly as the shear rate increases (Kang and

Pettitt, 1993). This shear-thinning property allows efficient pumping of xanthan polymer at high pumping (shear) rates. The mixing problem thus may be overcome by continuous medium circulation through the fibrous matrix and by rotating the fibrous matrix to enhance air, liquid, and cell contacts and to separate the xanthan polymer from the cells. In this new design, liquid media and air is passed through the porous fibrous matrix to ensure intimate contact with the immobilized cells. The centrifugal force generated form rotating the device is high enough to separate xanthan polymer from cells.

In this study, a novel, centrifugal, packed-bed bioreactor modified from the columnar bioreactor with spirally wound fibrous bed was employed for xanthan gum production. Two-stage repeated batch process was used to achieve high xanthan production rate. Enhanced oxygen transfer rate in the reactor was attained by applying centrifugal force and medium recirculation. 78

3.3 MATERIALS AND METHODS

3.3.1 Culture and Media

Xanthomonas campestris NRRL B-1459 was obtained from the Northern Regional

Research Laboratory (NRRL) of the U.S. Department of Agriculture (Peoria, IL). Stock cultures of X. campestris NRRL B-1459 were maintained on agar slants, which contained

10 g/L glucose, 3 g/L yeast extract, 5 g/L peptone, and 15 g/L agar, and were stored at

4°C. The culture was transferred once every two weeks to maintain good viability and stability for xanthan production.

Actively growing cells from a newly prepared slant culture (about 24-36 hour incubation time at 30°C) were loop inoculated into 500-mL Erlenmeyer flasks containing

100 mL liquid medium. After incubation for 24 hours at 30°C in an incubator-shaker, the

100-mL liquid culture was used to inoculate the fermentor containing 5 L of the medium.

Unless otherwise noted, the medium used in fermentation study consisted of 25 g/L glucose, 3 g/L yeast extract (USB, Cleveland OH), 2 g/L KjHPO^ 0.1 g/L MgSCV

7 H2O, and 500 ppm (v/v) antifoam A emulsion (Sigma Chemicals). Tap water was used in preparing the medium to provide trace elements. The medium was prepared in two parts: the first part contained all the basic medium components except for glucose, and the second part was a concentrated glucose solution. The initial medium pH was adjusted to 7 by adding 4N HC1. These two solutions were autoclaved at 121°C and 15 psig for 25 min„ allowed to cool to room temperature, and then mixed together aseptically. 79

3.3.2 Centrifugal* Packed-Bed Bioreactor (CPBR)

3.3.2.1 Bioreactor construction

The centrifugal, packed-bed bioreactor was constructed using a 5-L fermentor

(BioFlo II, New Brunswick Scientific Co., Edison, NJ) with the immobilized-cell matrix

and its support container attached to the central spinning shaft of the fermentor impeller

(Figure 3.1). The support container was a cylindrical cup (9 cm in diameter and 15 cm in height) with a central hollow core (2.7 cm in diameter), made of stainless plate with 3/16" round perforations on 1/4” staggered centers (Small Parts, Miami, FL). A selected fibrous sheet matrix (100% cotton towel; 5 mm x 15 cm x 80 cm) was overlaid with a crimped stainless steel wire cloth of the same dimension (Goodloe, Glitch Technology Corp.,

Dallas, TX), spirally wound around the vertical axis, and packed in the support container

(Figure 3.1(b)). The packed volume was -850 mL. Four tilted exterior baffles mounted on the wall of the container provided downloop flow pattern when the cup was rotated counterclockwise. A disk-blade turbine impeller installed at -1.5 cm below the cup provided turbulent mixing for the bottom part of the fluid in the fermentor. The centrifugal force generated from rotating the cup removed xanthan gum away from the immobilized- cell matrix to outer bulk broth (Figure 3.1(c)). The outer bottom part of the broth was pumped back into the center top of the rotating cup by a high-speed peristaltic pump

(Masterflex with No. 24 pump head, Cole-Parmer, Chicago, IL) at a flow rate of -270 mL/min. This recirculation stream was sprayed onto the rotating fibrous matrix through a brass nozzle (Nelson N-29C, L.R. Nelson Corp., Peoria, IL).

Sterile air was introduced into the medium through a ring sparger at the bottom of the fermentor vessel. Unless otherwise noted, the aeration rate was controlled at 1.5 vvm.

The temperature was controlled at 30±0.1°C. The reactor pH value was maintained at 80

Sparge Filter Motor I

pH /D O Meter Base ^ { Sample m : Circulation Pump

(a)

FIGURE 3,1: Schematic diagram of the centrifugal, packed-bed biorcactor: (a) reactor system. Spiral-rolled to create convoluted packing

Stainless steel screen topped with cotton terry cloth (b)

XV

>X'* ‘X -X X 0.* TeTc ’S• * i V * ’S • • S ♦ S v V 1

I

(c)

FIGURE 3.1: Schematic diagram of the centrifugal, packed-bed bioreactor: (b) configuration of the fibrous packing; and (c) fluid flow inside the bioreactor. 82

7.0±0.1 by adding 2.0 N potassium hydroxide. The dissolved oxygen tension in the broth was monitored by using a dissolved oxygen probe and analyzer (New Brunswick

Scientific, DO-40), which was connected to a chart recorder (LINSEIS L-4000 Digital

Flatbed Recorder) to record dissolved oxygen tension profiles.

3.3.2.2 Selection of the support matrix for cell immobilization

In order to select a suitable matrix for cell immobilization (adsorption), four types of fibrous matrices were investigated, including 100% cotton towel (with looping, ~5 mm thick), 100% cotton fabric sheet (without looping, ~1 mm thick), 50% cotton-50% polyester fabric sheet (-1 mm thick), and 100% polyester fabric sheet (~Im m thick). Each fibrous sample was tested for its cell adsorption capability by placing a small piece (3 cm x

3 cm) of the fibrous matrix in a fermentation flask containing 100 mL media. The flask was sterilized, inoculated with X. campestris, and incubated in an incubator-shaker at

30°C. The suspended cell density in each flask was monitored by measuring the optical density of the broth at regular time intervals for 50 hours. The attachment of X. campestris cells on each fibrous matrix was examined at the end of the experiment by scanning electron microscopy (SEM).

3.3.2.3 Bioreactor startup

The bioreactor was autoclaved twice at 121°C, 15 psig for 30 minutes and then filled with 5 liter sterile media containing 25 g/L glucose and 3 g/L yeast extract. The bioreactor was then inoculated with 100 mL flask culture. Unless otherwise noted, the cells in the bioreactor were first grown at pH 6 and 23°C, the optimal conditions for cell growth.

The bioreactor was aerated at a volumetric flow rate of 5.0 std. liter/min. During this period, the reactor impeller and the fibrous matrix were rotated at 150 rpm to provide good mixing. Then, when the reactor reached its highest cell density at -24 hours, the 83 fermentation conditions were changed to pH 7, 30°C, and 350 rpm to promote xanthan production and cell immobilization. The fermentation broth was then harvested at -50 hours when all glucose in the medium had been consumed and all cells had been immobilized onto the fibrous matrix. The reactor was ready for use in subsequent repealed batch fermentations.

3.3.2.4 Repeated batch fermentation

After the first batch fermentation, the bioreactor was operated as a repeated batch system that only the viscous broth was replaced with new sterile medium at the end of each batch fermentation as determined from the alkali addition rate approaching zero. New medium for next batch fermentation was prepared in advance and pumped into the bioreactor at a flow rate of -200 mL/min right after all xanthan broth had been pumped out from the bioreactor, which took about 25 minutes. The batch fermentation was repeated several times with 25 g/L glucose medium, and then with 50 g/L glucose medium to study the effect of C/N ratio on xanthan production.

The CPBR performance was studied under either liquid-continuous mode, where the fibrous bed was completely immersed in 5 L medium, or gas-continuous mode, where

90% of fibrous bed was exposed to air with only 2.5 L medium in the fermentor vessel.

The first study was conducted at liquid-continuous mode (5-liter medium) and 150 rpm rotational speed. The second study was also at the liquid-continuous mode but at 350 rpm rotational speed, which was the optimal rotational speed that provided high oxygen transfer rates without severe system vibration. The third study was conducted at gas-continuous mode (2.5 liters media) and at 350 rpm.

Samples of the fermentation broth were taken at proper time intervals. The cell density in the sample was determined immediately by measuring the optical density of the 84 cell suspension. Glucose and xanthan concentrations were determined after each biorcactor study was completed. The quality of xanthan produced from each batch was atso compared by measuring the apparent viscosity of the xanthan broth at a selected xanthan concentration of 18 g/L using the Brookfield viscometer.

3.3.2.5 Determination of immobilized cell density

The immobilized cell density in bioreactor was estimated at the end of each bioreactor study. First, all the liquid in the bioreactor was drained, and the liquid volume and OD were measured to estimate the total suspended cells (cells in the free solution) present in the reactor. In this study, however, no cells were found in the liquid medium.

The drained fibrous matrix was removed from the reactor and washed several times with water until almost all cells had been removed. The total volume of the washing water and its OD, as well as its dry weight, were measured and used to estimate the total amount of cells immobilized in the fibrous matrix.

3.3.2.6 Determination of cell viability

The relative viability of immobilized cells as compared to free cells was determined by a plate count method. Small pieces of Fibrous samples were cut off from the fibrous matrix and placed in test tubes containing sterile water. The immobilized cells in the fibrous matrix were then washed off from the matrix by vortexing for -3 minutes. The OD of the cell suspension was measured to estimate the total cell number in the sample. One mL of the cell suspension sample was then subjected to serial dilutions with saline before transferred onto agar plates. The total viable cell number was determined from the colony count (between 30-300) times the dilution of the sample. The same procedures were conducted with free cells grown in shake flasks for 24 hours. These free cells can be assumed to have 100% viability. A linear correlation between the plate count number and 85 OD reading was obtained from these cell samples. The total viable celt number for the

immobilized cell sample was then compared to the theoretical viable cell number (or total

cell number) at the measured OD value as obtained from the standard correlation. The ratio

between these two numbers was the relative cell viability for the immobilized cells.

3.3.2.7 Scanning electron microscopy

Several small pieces of the fibrous material were taken as samples from the drained

fibrous matrix. These samples were immersed in 2.5% glutaldehyde solution overnight,

rinsed with double distilled water completely, approximately 10 times, each time for 15

minutes. Then, they were gradually dehydrated with 20%-70% ethanol in the increment of

10% by holding them at each concentration for 30 min. The partially dehydrated samples

were left in 70% ethanol for overnight, then dehydrated progressively with 80% ethanol and twice with 95% and 100% ethanol for 30 minutes each time. These samples were then cryogenically dried at critical point with liquid CO 2. All steps, except that for the critical drying, were carried out at 4°C. The completely dried samples were coated with gold/palladium before taking SEM photograph using the JOEL model 820 SEM.

3.3.3 Analytical Methods

3.3.3.1 Cell density

Samples of fermentation broth were collected into centrifuge tubes. Depending on the broth viscosity, the broth samples were diluted with tap water by a factor of two to six, and the diluted solutions were centrifuged at 12,000 rpm (16,000xg) for 30 minutes at 5°C to precipitate the suspended cells. Cells were then resuspended in water and the optical density at 650 nm (ODeso) was measured using a spectrophotometer. The OD readings 86 were then compared to a standard correlation between OD and cell density (g/L). The cell density was proportional to OD when the optical density was below 0.5, with one unit of

OD equaling to 0.4 g/L cell. The total cell dry weight in the cell suspension was also determined after drying at 105°C for ~7 hours. The dry weight measurement was duplicated to reduce the experimental errors to within 0.2%.

3.3.3.2 Glucose concentration

The glucose concentration was determined by using a glucose analyzer (YSI Model

2700 SELECT, detection range: 0-25 g/L). Properly diluted cell-frce samples were presented to the needle port for automatic sample injection (adjustable from 5 to 65 microliters). The analysis is based on a biosensor membrane with immobilized glucose oxidase.

3.3.3.3 Xanthan gum concentration

Xanthan concentration in the fermentation broth was estimated from the broth viscosity with proper dilution (Shu and Yang, 1990). The viscosity of the cell-free fermentation broth was measured by using a Brookfield viscometer (RVTDII) with spindle

No. 1 at 100 rpm. The viscosity was then compared to a standard correlation between the viscosity and the xanthan concentration (g/L), which was linear when xanthan concentration was below 0.6 g/L. Samples were diluted with water to the proper concentration range before viscosity was measured. The final xanthan concentration of each batch fermentation was also determined by measuring the total dry weight of xanthan gum in the fermentation broth after purified by alcohol precipitation. The xanthan dry weight measurement was also used to verify the xanthan concentration determined from viscosity. 87

3.4 RESULTS AND DISCUSSION

3.4.1 Fibrous Matrix Selection

Among various types of fibrous matrices for cell immobilization studied, the 100% cotton towel with looping showed the fasted cell adsorption rate and had adsorbed almost all suspended cells by 50 hours of fermentation time (Figure 3.2). This fibrous matrix was thus chosen for use in the CPBR. The other matrices also adsorbed the cells, but at a much slower rate. In general, the hydrophilic cotton fiber was preferable to the hydrophobic polyester fiber. The scanning electron micrographs (Figure 3.3) also showed that more cells were attached to the 100% cotton towel than the other matrices. The rough surface of cotton fibers and the looping of the towel, in addition to its hydrophilicity, seemed to be important factors for cell adsorption to the surface. Also, the onset of cell adsorption seemed to coincide with the onset of stationary phase, where cell growth stopped and xanthan formation became significant. The production of xanthan gum probably helped cells to attach onto the fiber surface.

3.4.2 Reactor Startup

Typical kinetics of cell growth and xanthan production in the centrifugal, packed bed bioreactor during reactor startup is shown in Figure 3.4. Exponential cell growth but small amounts of xanthan biosynthesis occurred during the initial 24 hour period, which is also typical to xanthan gum batch fermentation with free cells. However, at -24 hours when the culture reached the stationary phase, the cell density in the fermentation broth rapidly decreased while xanthan production continued to increase, indicating that the cells were immobilized (adsorbed) onto the matrix. All suspended cells disappeared and were I E 3.2: RE U FIG

Cell Density (g/L) el dopin o aiu fbos arcs uig xanthan during matrices fibrous various to adsorption Cell 0.01 0.5 0 (plane sheet without looping); (c) 50% cotton + 50% polyester; and polyester; 50% + cotton 50% (c) looping); cotton 100% (b) without sheet looping); with (plane (towel 100%cotton (a) fermentation: d 100%polyester. (d) 10 ie (hr) Time 20 30 40 50 88 89

100% cotton (towel configuration) 100% cotton (plane configuration)

50% cotton + 50% polyester 100% polyester

■

I (0

FIGURE 3.3: Scanning electron micrograph of Xanthomonas campestris cells adsorbed on various fibrous matrices. 90

7.5 7.0 E a. 6.5 6.0 r g g ...... * * ...... *1 * ...... 0 10 20 30 40 50 60 70 80 90

32 30 28 a. 26 £ 2 4 . £ 22 ¥■■¥. W .ifl...... 0 10 20 30 40 50 60 70 80 90

e 400 I I I 11 V I I I I o ^ mid■"id' ' ' if "■ 'Ah' ' w E 300 s e- 200 O w 100 * “ ...... I....i 0 1 0 20 30 40 50 60 70 80 90

40 Cell £ 30 c Glucose o Xanthan «b e oat S 0.5 o

u Cell Density (g/L) J & j f c L U l L I llllllMII I 0 10 20 30 40 50 60 70 80 90

100 !■■■ 111 IP 80 ■ ■ 'W 60 _ D a H 40 m * D ■ ■ 20 - ■ a 0 0 10 20 30 40 50 60 70 80 90 Time (hr)

FIGURE 3.4: Typical kinetics of two-stage, repeated-batch xanthan gum fermentation in the CPBR. 91 immobilized onto the fibrous matrix by 50 hours, as indicated by the zero optical density of the broth. No free cells were detected in the subsequent repeated batch fermentation, indicating that all cells, both existing and from new growth, remained immobilized on the fibrous matrix. With all cells immobilized in the fibrous matrix, a cell-free xanthan broth was produced.

3.4.3 Liquid-Continuous Fermentation at 150 rpm

Figure 3.5 shows the kinetics for 4 repeated batch fermentations at the low rotational speed of 150 rpm. In the first two batches the xanthan production rate remained almost unchanged, whereas the following two batches with higher glucose-to-yeast extract ratio showed declined production rates. Figure 3.6 shows that both xanthan yield (Yp/S) and volumetric xanthan productivity (dP/dt) were lower for the last two batches. The reduced xanthan production at the later batches could be attributed to oxygen limitation in the system and possibly the reduced cell viability at the low rotational speed used in the fermentation. Agitation and aeration were relatively poor at 150 rpm when the xanthan concentration was higher than 1%.

3.4.4 Liquid-Continuous Fermentation at 350 rpm

To improve oxygen transfer and CPBR performance, the rotationat speed was increased to 350 rpm after cells were immobilized onto the matrix. As shown in Figure

3.7, the fermentation time required for each batch at this rotation speed was much shorter than that at 150 rpm. The fermentation time to reach -2.5% xanthan concentration was reduced to -40 hours, as compared to 50 hours or longer for conventional batch FIGURE 3.5: FIGURE

Concentration (g/L) 40 20 60 30 50 Kinetics of repeated-batch xanthan fermentation in CPBR at 150 rpm. 150 at CPBR in fermentation xanthan repeated-batch of Kinetics Cell 150rpm 50 Glucose 0 10 0 20 0 30 400 350 300 250 200 150 100 ie (hr) Time CP Xanthan

Cell Density (g/L) 92 Batch Number

FIGURE 3.6: Xanthan yield (Yp/s) and production rate (dP/dt) from the repeated- batch fermentation in CPBR at 150 rpm. 60 —T ■* i— 1— i— '— i— ■— i— •— i— '— i— i— r ->—i—»- 150 C 350 rpm rpm 50 - A Glucose A □ Xanthan $ s 40 - o • Cell □ * g 30 *5 / A □ B i V u 'a a J Q w 20 oC □ A U 10

01 i . ip. i . i 0 20 40 60 80 100 120 140 160 160 200 220 240 260 280 Time (hr)

FIGURE 3.7: Kinetics of repeated-batch xanthan fermentation in liquid-continuous CPBR (CPBR-LC).

•t*. 95 fermentation process. The productivity of the immobilized cells on the matrix was evaluated for long term operation by extending the repeated batch cycles. The reactor showed stable and consistent results for all six consecutive batches studied. As shown in

Figure 8, the xanthan yield remained at -85% and the quality of xanthan polymers produced, as determined by the solution viscosity of 1.8% xanthan concentration, remained unchanged for all the six batches. The reactor was contaminated by mold at the end of the last batch during medium change. The immobilized cells in the CPBR thus can be repeatedly used for xanthan production for a long term if the contamination problem can be avoided.

It is also noted that the volumetric xanthan productivity was significantly higher with the 5% glucose medium than with 2.5% glucose medium (Figure 3.8). The high C/N ratio in the 5% glucose medium was preferable for xanthan production. The productivity

(-0.7 g/L h) was also significantly higher than that of the conventional batch fermentation process (-0.5 g/L h). It is clear that cell immobilization using hydrophilic fibrous matrix as the carrier can produce cell-free xanthan broth at high production rate. After the cells were successfully immobilized, they became available for reuse in subsequent batches, thus eliminating the long period for call growth required in the conventional batch fermentation.

The reactor had high, stable productivity and can be efficiently operated at repeated batch mode.

3.4.5 Gas-Continuous Fermentation

Since all cells were immobilized in the fibrous matrix, xanthan gum production could only take place when the glucose medium was in contact with the fibrous matrix. I E 3.8: RE U FIG □ >- Viscosity (cP) 'a, 10 ■ 3100 2900 3300 Xanthan yield (Yp/s), poduction rate (dP/dt), and quality (in terms of terms (in quality and (dP/dt), rate poduction (Yp/s), yield Xanthan 0.0 0.4 0.2 0.8 0.6 1.0 --- the apparent viscosity of 1.8% xanthan solution) from the repeated- repeated- the from solution) xanthan 1.8% CPBR-LC. in fermentation of batch viscosity apparent the ------.% ata a ser ae s'1 ' s 6 rate shear at Xanthan 1.6% 2 4 6 5 4 3 2 1 1 1 q ------0 0 . Gu 03 E50 l/. YE Glu/0.3 5.0 YE 0.3 / Glu 2.5 1 1 ------ac Number Batch q 1 1 o 9 ------^ ------1 1 ------1 1 ------•— 1 0.0 0.2 . fc. 0.6 0.4 0.4 0.8 1.0 *o £ t JS ■o

96 97

The fermentation time thus should be further reduced by increasing the liquid-cell contact or reducing the liquid volume in the reactor vessel.

Figure 3.9 shows the fermentation results from gas-continuous operation with 2.5 liters media in the fermentor vessel. As can be seen that the fermentation time for each batch cycle was dramatically reduced to 24-26 hours, half of the time required for the conventional process. The xanthan yield, volumetric xanthan productivity, and xanthan quality from eight batches are shown in Figure 3.10. The volumetric productivity, based on the total liquid volume in the fermentor vessel, was ~1 g/L h. The reactor productivity was -2.5 g/L h if only the actual reactor working volume (the volume occupied by the rotating fibrous matrix) was used in the calculation. The productivity can be further increased by increasing the medium recirculation rate, which increases gas-liquid and cell- liquid contacts in the fibrous matrix. The oxygen transfer rate should be enhanced by increasing liquid recirculation rate. This would result in a higher xanthan production rate.

As can be seen from Figure 3.10, both the xanthan yield and volumetric xanthan productivity were about the same for fermentations with 5.0% glucose/0.3% yeast extract and 2,5% glucose/0.15% yeast extract. Thus the enhanced xanthan production rale at the higher C/N ratio also was achieved at a relatively low yeast extract concentration. This clearly showed that C/N ratio is the major factor in affecting the xanthan gum production rate by the immobilized cells.

3.4.6 Cell Immobilization

Fibrous samples from various parts of the fibrous matrix were examined for cell density distribution at the end of each reactor study. It was found that cells were evenly FIGURE 3.9: FIGURE

Concentration (g/L) Kinetics of repeated-batch xanthan fermentation in gas-continuous CPBR(CPBR-GC); liquid 40 60 20 30 50 10 0 volume wasreduced from 5.0 L toL 2.5 after first batch. 0 < i a 0 0 0 O 0 10 4 10 6 20 2 240 220 200 160 160 140 120 100 SO 60 40 20 . L 5.0 Xanthan □ Cell • Glucose * . L 2.5 ie (hr) Time u p ■M n 4i ft* S 60 J 30 i E 3.10: RE U FiG

Viscosity (cP) Y p/s, 2900 3100 3300 Xanthan yield (Yp/s), production rate (dP/dt), and quality (in terms of of terms (in quality and (dP/dt), rate production (Yp/s), yield Xanthan their apparent viscosity of 1.8% xanthan solution) from the repeated- repeated- the from solution) CPBR-GC. xanthan in 1.8% fermentation batch of viscosity apparent their __ o - Xanthan 1.6% ■ —1 L. ------8 i 1 ------__ 4 6 5 4 3 8 i 1— __ ac Number Batch i 1 T —i— __ a i __ A a o

t ha rt 6s'1. ' s 6 rate shear at 1 _...... 1 — o - o 0 1 ------_____ 1 1 __ --- 99 1(H) distributed onto the fibrous matrix. Figure 3.11 shows the attachment of Xanthomonas campestris on the fiber surface. As can be seen in Figure 3.11(a), the towel configuration of the matrix was able to provide relatively large surface for cell attachment. The immobilized cells were well spreaded on the fiber surface (Figure 3.11(b)). They adsorbed on the surface as individual cells and without the formation of large lumpy colonies (Figure

3.11(c) and (d)).

At the end of each reactor study, the amount of immobilized cells in the bioreactor was found to be 34 g for the liquid continuous reactor and 38.4 g for the gas-continuous reactor. This gave an equivalent liquid cell density of 6.8 g/L and 15.36 g/L, respectively, which were at least three to seven times of that obtained in the conventional free-cell batch xanthan fermentation (< -2 g/L). It is clear that the higher immobilized cell density also contributed to the higher productivity for CPBR. However, the increase in CPBR productivity was less proportional to the increase in the total cell density, indicating that cell viability in the CPBR might be significantly lower than 100%. The cell viability of the immobilized cells in the CPBR at the end of the study was only 60% of that for free suspended cells. This explained partially why the volumetric xanthan productivity of CPBR was not proportional to the total cell density in the bioreactor.

3.4.7 Free Cell Batch Fermentation in STR

For comparison purpose, batch xanthan fermentations with free cells in conventional stirred tank bioreactor (Marubishi MD-300) were also studied. Kinetics of cell growth, xanthan production, and glucose consumption for two batch xanthan fermentations at 30°C and pH 7.0 are shown in Figures 3.12 and 3.13, respectively. At the early stage of fermentation, the inoculated culture showed exponential growth. The 101

«> ■ g crLayer

Inner Layer

FIGURE 3.11: Scanning electron micrographs: (a) the configuration of towel matrix; (b) the attachment of Xanthomonas campestris cells on the fiber surfaces; (c) matrix sample from outer packing layer; and (d) matrix sample from inner packing layer. 102

30

mi Cell "Sb 20 1 Glucose s © ° 8 0 e ‘■S □ l/i « Xanthan c +*b e £ s 10 e V o o V

20 40 60 ao 100 £ eo H O O 0 20 40 60 80 0.3 0.6 0.6 0.2 OTR 0.4 ii 0.1 0.2 Qfi dP/dt/Xs T! M H o.o 0.0 O o 20 40 60 8U

20 40 60 Time (h)

FIGURE 3,12: Kinetics of batch xanthan fermentations in stirred tank reactor (STR) with 2.5% glucose. u mi cji a cci i v i n v p m m|iii111 i i l >• □ o □ o> m o Cell Density (g/L) i...... i....Q.Jl b tr o oi o ui o> ° o o o o m o -1 a ro ro 4 ___DOT (%) x Concentration (g/L) ^ oooooo o o o o o o o ro o ro o

o ro o u o (g/g Cell/h) o o o — o ro dP/dt/X s b> b OTR (g/L h) CD o o ro o O) o o o u Op o p o o

o a t a o o o o o (mg/gCell/h) o o o o o o Specific OUR a> o ro o o o o o io o H 3 rp

FIGURE 3.13: Kinetics of batch xanthan fermentations in stirred tank reactor (STR) with 5.0 % glucose. 104 culture then reached the stationary phase. It took -24 hours for cells to reach a maximum cell density with the low glucose (e.g. 2.5%) medium. Large amount of xanthan formation

occurred when cell growth stopped. For the high glucose medium (5.0%), cell growth was much slower and was in parallel with xanthan formation for most of the fermentation time. The fermentation time was long - -70 hours for the 2.5% glucose medium and -120

hours for the 5.0% glucose medium.

Dissolved oxygen tension (DOT) monitored in the medium decreased as the fermentation progressed, dropping gradually from 100% to ca. 50%. As the broth became more viscous at 2% - 3% xanthan, DOT dropped quickly down to below 10%. The DOT needs to be maintained above 20% to prevent any adverse effect caused by oxygen on xanthan limitation (Moraine and Rogovin, 1973; Flores et a i, 1994). The oxygen transfer rate (OTR) during the batch fermentation was estimated from DOT and the values for kLa and C* determined in Chapter IV. As also shown in these two figures, OTR decreased as xanthan concentration increased. The specific xanthan productivity also seems to be in parallel with the specific oxygen uptake rate (OUR).

Both the cell yield and specific growth rate of X. campestris in the stirred tank reactor (STR) were found to decrease with increasing C/N ratio or the glucose to yeast extract ratio in the medium (DeVuyst et a i , 1987; Lo, 1993). On the other hand, the xanthan yield and specific xanthan production rate increased with increasing the glucose to yeast extract ratio in the medium. 105

3.4.8 Comparison between STR and CPBR

Figure 3.14 shows the comparison of volumetric xanthan productivities (dP/dt) and

specific xanthan productivities (dP/dt/Xs) from STR, CPBR-LC, and CPBR-GC. As can

be seen in Figure 14(a), CPBR gave higher volumetric productivity than conventional

STR, mainly because of its higher cell deasity in the reactor. However, CPBR had a lower specific productivity (Figure 3.14(b)) because of the relatively low cell viability. This

indicated that the oxygen transfer rate in the CPBR needs to be further increased in order to fully utilize the high density of immobilized cells in the fibrous matrix. Detailed discussion of OTR is given in Chapter IV,

As shown in Figure 3.15, the specific xanthan productivity is a function of specific oxygen uptake rate (OUR) during xanthan gum fermentation. The performance of CPBR can be further improved to a higher level than that was reached in this study by increasing the liquid recirculation rate and increasing the rotational speed of the fibrous matrix.

3.5 CONCLUSION AND RECOMMENDATION

Since xanthan gum is produced as a secondary metabolite and its production can be non-growth associated, the optimal conditions for cell growth and xanthan biosynthesis are quite different. In general, optimal cell growth requires a relatively low temperature (22-

24°C) (Shu and Yang, 1990), low pH (6) (Thonart et al., 1985), high oxygen demand

(182 mg 0 2 /g cell/hr) (Herbst et at., 1987), and low medium C/N ratio (DeVuyst et al .,

1987; Lo, 1993), whereas higher temperature (31-33°C) and pH (7), lower oxygen demand (148 mg 0 2 /g cell/hr), and higher C/N ratio are needed for optimal xanthan production. Therefore, separation of cell growth and xanthan production into two stages 106

2.5 liter 2.5/0.3 1.0 5.0/0.3

0 8 5 liter ■85 0.6 5 liter

ISu -O 0.4 s _3 0.2 > 0.0 STR CPBR-LC (a)

0.3 5 liter 2.5/0.3 5.Q/0.3 ■« "t: > ** u 0.2 9 CD

5 liter Pcu w IS 2.5 liter

0.0 STR CPBR-LC CPBR-GC

(b)

FIGURE 3.14: Comparison of (a) volumetric xanthan productivity; and (b) specific xanthan productivity of batch fermentations in STR, CPBR-LC, and CPBR-GC. IUE : 5 J 3 FIGURE

dP/dt/Xs (g/g Cell/h) 0.0 0.3 0.4 0.1 0.2 0.5 0.6 i » i ii r i i Specific xanthan productivity as a function of specific oxygen oxygen specific of function a as productivity xanthan Specific uptake rate during xanthan gum fermentation. gum xanthan during rate uptake < 0°^ ° r

» (CPBR) O i i ' i—*—i i' ■ ii t ■ i i _ ' . i I ■ I ■ i ■ . I ' i (STR) : o 500 107 108 improved xanthan fermentation with enhanced xanthan production rate and yield as demonstrated in this study.

The new centrifugal, packed-bed bioreactor was able to produce xanthan gum at a productivity twice of that for the present industrial process. The intimate air, liquid, and cell contacts achieved via passing liquid medium and air through the porous fibrous matrix achieved an enhanced oxygen transfer rate and high xanthan productivity. The cells immobilized in the bioreactor was able to be repeatedly used for xanthan fermentation to achieve continuous production of cell-free xanthan broth. Therefore, subsequent cell- removal from fermentation broth is eliminated. The production of cell-free fermentation broth could further benefit the downstream ultrafiltration process by eliminating membrane fouling caused by cells and cell debris (DNA, RNA, etc.) that would otherwise be present in the xanthan broth. The immobilized cell density in the reactor can be increased to further increase the reactor productivity. However, this would require a parallel increase in oxygen transfer rate.

3.6 REFERENCES

Atkinson, B. 1986. Immobilized cells, their applications and potential. Ch. 1. In: C. Webb, G. M. Black, and B. Atkinson (eds.), Process engineering aspects of immobilized cell systems. Pergamon Press, New York.

Atkinson, B., Cunningham, J. D., Pinches, A. 1984. The biomass hold-ups and overall rates of substrate (glucose) uptake of support particles containing a mixed microbial culture. Chem. Eng. Res. Dev. 62: 155-159.

Behrens, U., Klima, M., Fiedler. S. 1980. Growth and accumulation of polysaccharide by Xanthomonas campestris. Z. Allg. Mikrobiol. 20: 209-213.

Black, G. M. 1986. Characteristics and performance of immobilized cell reactors, Ch. 5. In: C. Webb, G. M. Black, and B. Atkinson (eds.), Process engineering aspects of immobilized cell systems. Pergamon Press, New York. 109 Cottrell, I. W., Kang, K. S., Kovacs, P. 1980. Xanthan gum, Ch. 24. In: R. L. Davidson (ed.), Handbook of water-soluble gums and resins. McGraw-Hill, New York.

DeVuyst, L., VanLoo, J., Vandamme, E. J. 1987. Two-step fermentation process for improved xanthan production by Xanthomoms campestris NRRL B-1459. J. Chem. Tech. Biotechnol. 39: 263-273.

DiCosmo, F., Tanaka, H., Neumann, A. W. 1994. Cell immobilization by adsorption to glass fibre mats. In: I. A. Veliky and R. J. C. McLean (eds.), Immobilized biosystems: theory and practical applications. Blackie Academic & Professional, Glasgrow, UK.

Flores, F., Torres, L. G., Galindo, E. 1994. Effect of the dissolved oxygen tension during cultivation of X. campestris on the production and quality of xanthan gum. J. Botechnol. 34: 165-173.

Foerberg, C., Enfors, S. O., Haeggstroem, L. 1983. Control of immobilized, nongrowing cells for continuous production of metabolites. Eur. J. Appl. Microbiol. Biotechnol. 17: 143-151.

Fowler R„ Balasundaram, V. 1989. De-aeration under the action of centrifugal force. In: Proceedings of water management offshore, Aberdeen, Scotland.

Funahashi, H., Maehara, M., Taguchi, H„ Yoshida, T. 1987. Effects of agitation by fiat- bladed turbine impeller on microbial production of xanthan gum. J. Chem. Eng. Jpn. 20: 16-22.

Galindo, E., Nienow, A. W. 1992. Mixing of highly viscous simulated xanthan fermentation broths with the Lightnin A-315 impeller. Biotechnol. Prog. 8:233-239.

Galindo, E., Nienow, A. W. 1993. Performance of the Scaba 6SRGT agitator in mixing of simulated xanthan gum broths. Chem. Eng. Technol. 16: 102-108.

Herbst, H., Peters, H. U., Suh, I. S., Schumpe, A., Deckwer, W. D. 1987. Oxygen transfer during xanthan fermentation. 1987 AIChE Annual Meeting, 175a, New York.

Herbst, H., Schumpe, A., Deckwer, W. D. 1992. Xanthan production in stirred tank fermenlors: oxygen transfer and scale-up. Chem. Eng. Technol. 15: 425-434.

Himmelsbach, W. 1985. Studies on mass and heat transfer and on homogenization in stirred fermentors using xanthan as an example. Chem. Ing. Tech. 57; 548-549.

Jcanes, A. R„ Pittsley, J. E., Senli, F. R. 1961. Polysaccharide B-1459: hydrocolloid polyelectrolytes produced from glucose by bacterial fermentation. J. Appl. Polymer Sci. 5: 519-526.

Ju, L. K., Zhao, S. 1993. Xanthan fermentation in water/oil dispersions. Biotechnol. Techniq. 7: 463-468.

Kawase, Y., Tsujimura, M. 1994, Enhancement of oxygen transfer in highly viscous non- Newtonian fermentation broths. Biotechnol. Bioeng. 44: 1115-1121. 110 Kennedy, J. F.; Bradshaw, I. J. 1984. Production, properties and applications of xanthan, pp. 319-371. In: M. E. Bushel) (ed.), Progress in industrial microbiology, vol. 19. Elsevier Science, New York.

Kessler, W. R., Popovic, M. K., Robinson, C. W. 1993. Xanthan production in an extemal-circulation-loop airlift fermenter. The Can. J. Chem. Eng. 71: 101-106.

Lewis, V. P., Yang S. T. 1992. Continuous propionic acid fermentation by immobilized Propionibacterium acidipropionici in a novel packed-bed bioreactor. Biotechnol. Bioeng. 40: 465-474.

Lo, Y. M. 1993. Effects of medium composition on product formation and cell growth in xanthan gum fermentation. M.S. thesis, The Ohio State University, Columbus, OH.

Long, R„ Roubicek, R. 1988. Centrifugal film reactor - a new concept for cell cultivation. Bioprocess Eng. 3: 73-77.

Margaritis, A., Pace, G. W. 1985. Microbial polysaccharides, pp. 1005-1043. In: M. Moo-Young (ed.), Comprehensive Biotechnology, vol. 3. Pergamon Press, New York.

Misra, T. K., Barnett, S. M. 1987. Evaluation of a novel foam fermentor in the production of xanthan gum, pp. 227-237. In: C. S. Ho and J. Y. Oldshue (eds.), Biotechnol. Processes. AIChE, New York.

Mohr, R. J., Khan, A. S. 1987. Higee - a new approach in groundwater clean-up. In: Proceedings of the AQTE Conference, Montreal, Canada.

Moraine, R. A., Rogovin, P. 1973. Kinetics of the xanthan fermentation. Biotechnol. Bioeng. 15: 225-237.

Moraine, R. A., Rogovin, P. 1971. Xanthan biopolymer production at increased concentration by pH control. Biotechnol. Bioeng. 13: 381-391.

Nakajima, S., Funahashi, H., Yoshida, T. 1990. Xanthan gum production in a fermentor with twin impellers. J. Ferment. Bioeng. 70: 392-397.

Nienow, A. W. 1984. Mixing studies on high viscosity fermentation processes xanthan gums. The World Biotech Report 1: 293-304.

Pace, G. W„ Righelato, R. C. 1980. Production of extracellular microbial polysaccharides, pp 41-70. In: A. Fiechter (ed.). Advances in biochemical engineering, vol. 15. Springer-Verlag, New York.

Parts, A. G., Elbing, E. 1975. The polymerisation of acrylonitrile in centrifugal fields. Paper presented at American Chemical Society Meeting, Philadelphia. Polym. Prepr. 16: 211-216.

Peters, H. U., Herbst, H., Hesselink, P. G. M„ Liinsdorf, H., Schumpe, A., Deckwer, W. D. 1989. The influence of agitation rate on xanthan production by Xanthomonas campestris. Biotechnol. Bioeng. 34: 1393-1397. I l l Peters, H. U., Suh, I. S., Schumpe, A., Deckwer, W. D. 1992. Modeling of batchwise xanthan production. The Can. J. Chem. Eng. 70: 742-750.

Pinches, A., Patient, L. J. 1986. Rate and yield relationships in the production of xanthan gum by batch fermentation using complex and chemically defined media. Biotechnol. Bioeng. 28: 1484-1496.

Pons, A., Dussap, C. G., Gros, J. B. 1989. Modelling Xanthotnonas campestris batch fermentation in a bubble column. Biotechnol. Bioeng. 33: 394-405.

Qadeer, M. A., Baig, S. 1987. Microbial synthesis of xanthan gum by repeated-batch process. Pakistan J. Sci. Ind. Res. 30: 886-889.

Robinson, D. K., Wang, D. I. C. 1988. A transport controlled bioreactor for the simultaneous production and concentration of xanthan gum. Biotechnol. Prog. 4: 231-241.

Roubicck, R„ Feres, V. 1987. Equipment for enhanced mass transfer in chemical and biochemical processes. U.S. Patent No. 4,657,677.

Schumpe, A., Diedrichs, S., Hesselink, P. G. M., Nene, S., Deckwer, W. D. 1991. Xanthan production in emulsions, pp. 196-199. In: Proceedings of International Symposium on Biochemical Engineering - Stuttgart, 2nd.

Setford, S. J., Barker, P. E„ Ganetsos, G. 1994. Combined bioreaction and separation in centrifugal fields. Bioseparation 4: 117-126.

Shu, C. H„ Yang, S. T. 1990. Effect of temperature on cell growth and xanthan production in batch cultures of Xanthomonas campestris. Biotechnol. Bioeng. 35: 454- 468.

Smelser, S. C., Woodcock, K. E., Meyer, H. S., Fowler, R. 1990. Selective acid gas removal using the Higee absorber. Proceedings of AIChE Meeting, Orlando, FL.

Suh, I. S., Schumpe, A., Deckwer, W. D. 1991. Xanthan production in bubble column and air-lift reactors. Biotechnol. Bioeng. 39: 85-94.

Tecanle, A., Choplin, L. 1993. Gas-liquid mass transfer in non-Newtonian fluids in a tank stirred with a helical ribbon screw impeller. The Can. J. Chem. Eng. 71: 859-865.

Thonart, Ph., Paquot, M., Hermans, L., Alaoui, H. 1985. Xanthan production by Xanthomonas campestris NRRL B-1459 and interfacial approach by zela potential measurement. Enzyme Microb. Technol. 7: 235-238.

Yang, S. T„ Zhu, H,, Li, Y. 1994. Continuous propionate production from whey permeate using a novel fibrous bed bioreactor. Biotechnol. Bioeng. 43: 1124-1130.

Zaidi, A., Ghosh, P., Schumpe, A., Deckwer, W. D. 1991. Xanthan production in a plunging jet reactor. Appl. Microbiol. Biotechnol. 35: 330-333.

Zhu, H. 1995. A novel fibrous-bed bioreactor for mammalian cell culture. Ph.D. dissertation, The Ohio State University, Columbus, OH. CHAPTER IV

OXYGEN TRANSFER IN VISCOUS XANTHAN BROTH

4.1 SUMMARY

The volumetric mass transfer coefficient, kja. and oxygen transfer in the viscous xanthan fermentation broth for a centrifugal, packed-bed bioreactor (CPBR) was studied.

Other bioreactor systems, including the conventional stirred-tank bioreactor with disc turbine (STR-DT) or marine propeller (STR-MP), and the water-in-oil emulsion system

(STR-WIO), were also studied. STR-WIO showed the highest maximal kta values among those systems studied, while STR-MP was found to be the poorest. However, practical difficulties exist in recovering xanthan polymers from such an emulsion system. In CPBR it was found that, at xanthan concentrations lower than 2%, the maximal kia values were slightly lower than those of STR-DT. On the contrary, at xanthan concentrations higher than 2%, the maximal k^a values in CPBR were higher than those of STR-DT and close to the values of STR-WIO system, indicating improved oxygen transfer character in CPBR at high xanthan concentrations. The recirculation stream used to provide better mixing and gas-liquid contact in CPBR had an important role in providing aeration for the gas- continuous CPBR system.

112 m

Keywords: xanthan fermentation, oxygen transfer, aeration, agitation, bioreactor.

4,2 INTRODUCTION

Xanthan gum, a microbial polysaccharide, is commercially produced as a fermentation product by Xanthomonas campestris (Margaritis and Pace, 1985). The worldwide acceptance and applications of xanthan gum mainly take advantage of its excellent and unique rheological properties as well as its high stability at extremes of pH and shear, and high resistance to thermal degradation (McNeely and Kang, t973; Moraine and Rogovin, 1973; Behrens e ta i, 1980; Pinches and Pattern, 1986; Peters et al., 1992).

High viscosity xanthan solutions can be formed even at a low concentration in either hot or cold water (Jeanes, 1974; Cottrell e ta i, 1980).

However, the conventional batch fermentation process for xanthan gum manufacturing using stirred-tank reactors (STR) suffers from the high viscosity simultaneously occurred with increasing xanthan concentration, resulting in oxygen limitation (Sittig, 1983). High solution viscosity deteriorates the mass transfer, mixing, and heat transfer characteristics of the fermentor (Peters e ta i, 1989; Gonzales et a i ,

1989), as well as the pumping capacity of the impeller (Solomon e ta i, 1981). Therefore, oxygen transfer becomes problematic in viscous xanthan fermentation and the maximum xanthan concentration achieved in STR is usually limited to 3.0% w/v (Pace and Righelato,

1980; Cottrell e ta i, 1980; Solomon e ta i, 1981; Lim e ta i, 1984; Herbst e ta i, 1989). In general, oxygen limitation occurred in a xanthan fermentation when the DOT is below 20%)

(Flores et a i, 1994; Moraine and Rogovin, 1973). Furthermore, the specific rates and yields of xanthan production, as well as xanthan quality were found to be dependent on the dissolved oxygen tension, and more specifically on the volumetric oxygen transfer rate 114 (OTR, in g O 2/L h) in the fermentor (Holzwarth, 1978; Mukhopadhyay and Das, 1994). A

linear relationship between specific oxygen uptake rate (OUR, in mg 0 2 /g cell/h) and the specific growth rate, /i, was also reported (Pinches and Patient, 1986). According to

Moraine and Rogovin (1973), the specific xanthan production rate drops from almost 0.5 g/g cell/h initially to less than 0.1 g/g cell/h at the end of a batch fermentation. The drop in productivity can be attributed to the decrease in oxygen transfer rate caused by the increased viscosity and inappropriate agitation at high viscosity.

The provision of oxygen for a fermentation cannot be estimated simply by overall demand, because the metabolism of the culture is affected by the concentration of dissolved oxygen in the broth. OTR has long been used to indicate the aeration ability of a fermentation system in which various aeration requirement can be performed. The OTR from air bubble to the liquid phase is described by the equation;

OTR = ^ = kLa(C*-Ct) (4.1) at where CL (mmoles dm*3) is the concentration of dissolved oxygen in the fermentation dC broth; / (hour) is time; —— is the change in oxygen concentration over a time period, i.e. dt the oxygen transfer rate (OTR, mmole 0 2 dm*3 h*1); kL (cm h*1) is the mass transfer coefficient; a (cm2 cm*3) is the gas/liquid interface area per liquid volume; and C* (mmole dm*3) is the saturated dissolved oxygen concentration. Practical difficulty exists in measuring both kL and a separately in a fermentation; therefore, the two terms are generally combined in the term kLa (h*1), the volumetric transfer coefficient. The value of kta is a measure of the aeration efficiency and capacity of a fermentor under the test conditions, the larger the kLa the higher the aeration capacity of the system. kLa is thus frequently used as an indication of the oxygen transfer rate that can be achieved in a bioreactor. The determination of k,a of a fermentor becomes essential in order to establish the aeration 115 efficiency of a system and to quantify the effects of operating variables on the provision of

oxygen.

A number of factors have been shown to affect the kLa value in a fermentation

system. Such factors include the air-flow rate employed, the degree of agitation, the

rheological properties of the culture broth, and the presence of antifoam agents. The air­

flow rate in the range of 0.5 to 1.5 vvm has a relatively small effect on kLa value in

conventional agitated systems. The kLa in xanthan broth decreases mainly from the effect

of viscosity on interfacial area, a.

The rheology of the fermentation broth has a profound effect on the oxygen transfer

efficiency of an agitated fermentor (Funahashi et a i, 1987; Peters et al., 1989). The effective viscosity for a pseudoplastic fermentation broth increases with decreasing the

shear rate following the Osterwald-deWaele law:

V = (4.2)

The viscosity of xanthan fermentation broth increases dramatically in the final phase of

fermentation (Herbst e ta i, 1987). The influence of the effective viscosity (10-1000 cP) on

the kLa in STR was observed to be:

kLa^u~°J (4 .3)

Mechanical agitation in a fermentor is generally required to achieve good heat and

mass transfer for viscous fermentations (Margaritis and Zajic, 1978; Pace and Righelato,

1980; Bhavaraju et a i, 1978; Henzler, 1981). The design of the agitation system for use in polysaccharide fermentations requires special attention to give the correct distribution of power to ensure good culture homogeneity and turbulence, to minimize bubble coalescence, to promote small bubble formation, and to achieve adequate fluid movement al 116 heat transfer surfaces (Sittig, 1982; Margaritis and Pace, 1985; Galindo and Nienow,

1992).

Various approaches towards the estimation of kia in a STR with viscous pseudoplastic fluid have been studied and were later summarized by Herbst et a i (1992).

Considerable research attention has been given to the oxygen transfer characteristics of

STRs with various types of impellers because of the widespread use of STRs in the fermentation industry (Ahmad et a/., 1994). Recently, extensive efforts have also been made to develop new types of bioreactors and agitation designs with improved aeration and oxygen transfer rate and lower energy costs (Nienow, 1984; Himmelsbach, 1985; Herbst et al.* 1987; Misra and Barnett, 1987; Pons etai., 1989; Nakajima e ta i, 1990; Suh et a i,

1991; Zaidi e ta i, 1991; Galindo and Nienow, 1992; Herbst et a i, 1992; Peters et a i,

1992; Galindo and Nienow, 1993; Kessler e ta i , 1993; Tecante and Choplin, 1993;

Kawase and Tsujimura, 1994). Also, high OTR and good mixing characteristics have been reported for water-in-oil emulsion systems used in viscous xanthan fermentation (Schumpe et a i, 1991; Ju and Zhao, 1993). Due to the high oxygen diffusion coefficient and solubility in oil, which were determined to be 2.49 x 10 5 cm2/s and 10.635 x 103 mol

0 2 /L-atm, respectively, as compared to 2.11 x 10'5 cm2/s and 1.298 x 10-3 mol O^/L-atm in water at 22"C and ! atm (Ju and Ho, 1989), the water-in-oil emulsion system was reported to provide better oxygen transfer in the agitated fermentor system.

Recently, a novel centrifugal, packed-bed immobilized cell bioreactor (CPBR) was developed in our research group to improve xanthan gum fermentation (Yang e ta i, 1994;

Lo et a i, 1995). This new bioreactor was inspired from the Higee concept, a device that uses centrifugal force to provide high mass transfer in a wide variety of process applications (Mohr and Khan, 1987; Fowler and Balasundaram, 1989; Smelser e ta i,

1990). The most significant characteristic of CPBR is the high viable cell density achieved 117 in the matrix via natural attachment for cell immobilization (Chapter III). In order to achieve higher productivity, higher mass (especially for oxygen) transfer rates in the system need to be maintained. In CPBR, liquid media and air were passed through the porous fibrous matrix to ensure intimate contact with the immobilized cells. The aeration can be improved by continuous medium circulation through the fibrous matrix and by rotating the fibrous matrix to enhance contacts and to separate the xanthan polymer from the cells. Enhanced oxygen transfer and conditions for improved mixing in a centrifugal film bioreactor, particularly of very viscous liquid media, have also been described by several authors

(Roubicek and Feres, 1987; Long and Roubicek, 1988). The use of centrifugal bioreactors, developed for bioseparation, also has been the subject of recent studies for applications in enzyme and fermentation technologies (Svcnsson et a l, 1957; Parts and

Elljing, 1975; Setford e ta i, 1994). Investigations into the oxygen transfer efficiency of such a system is essential to evaluate and to improve its performance during highly viscous xanthan fermentation (Lo, 1993; Lo, 1995).

The objective of this study was to evaluate the aeration efficiency of CPBR and several other bioreactor systems, including STR with disc turbine (STR-DT), STR with marine propeller (STR-MP), and STR with water-in-oil emulsion (STR-WIO). The effects of agitation speed and xanthan concentration on the system mass transfer characteristics, including the volumetric mass transfer coefficient ( kua ) and oxygen transfer rate (OTR), were studied. Oxygen transfer efficiency as measured by the kua value for the various bioreactor systems was compared and discussed in this paper. 118

4.3 MATERIALS AND METHODS

4.3.1 Xanthan Fermentation Broth

Xanthan fermentation broths with concentrations of ca. 20-35 g/L (wt/v) were

produced from a glucose medium at pH 7.0 by batch fermentation using Xanthomonas campestris NRRL B-1459 following previously described procedures (Lo, 1993).

Solutions with lower concentrations were obtained via additional dilution of original xanthan broths with water.

4.3.2 Stirred Tank Reactor (STR)

A 5-L bench-top fermentor (Mirubishi MD-300) was used as the STR. Two different types of impellers were studied for their effects on kya in xanthan fermentation broth. The 6-blade disc turbine impeller (STR-DT), which is currently used in xanthan fermentation, produces radical flow pattern in mixing the fluid (Figure 4.1(a)). Another type of impeller widely used in industry is the three-blade marine propeller (STR-MP), which causes axial flow pattern in mixing the fluid in a baffled tank (Figure 4.1(b)). Air was introduced into the broth through the ring sparger at the bottom of the reactor vessel.

Unless otherwise noted, the reactor contained 5 liter xanthan fermentation broth was aerated at 5 L air/min, and was maintained at 30°C.

The same STR reactor with disc turbine impeller was also used in studying the water-in-oil emulsion (STR-WIO, Figure 4.1(c)). n-Hexadecane was employed as the oil phase to form a 50/50 water-in-oil emulsion. J—, oil « * ■ « •■ / .* • ^ • * * r • L I I • *-»)\

(a) (c)

(b) <

FIGURE 4.1: Schematic diagrams and flow patterns of bioreactors studied: (a) STR-DT; (b) STR-MP; (c) STR-WIO; and (d) CPBR. 120 4.3.3 CPBR

The CPBR was constructed inside a 5-L BioFlo II fermentor (New Brunswick

Scientific Co., Edison, NJ). A schematic diagram of the bioreactor system is shown in

Figure 4.1(d). In order to provide better air-contact and mixing in the reactor, a circulation stream was installed to bring the outer bottom broth back into the top center of the centrifugal cup and then nozzle-sprayed onto the matrix via the aid of a high-speed peristaltic pump (Masterflex with No. 13 pump head, Cole-Parmer, Chicago, IL). Two types of operation to investigate the OTR inside the biorcaetor were conducted: the liquid- continuous mode (CPBR-LC), in which the whole matrix was immersed in 5 liters broth; and the gas-continuous mode (CPBR-GC), in which the matrix was kept above the broth

(2.5 liters). A recirculation stream was used to provide better mixing and gas-liquid contact in the rotating fibrous matrix. The effect of recirculation on kia was studied with an effective recirculation rate of 270 mUmin. Air was provided form the ring sparger at the bottom of the reactor, at a volumetric rate of 5 liters/min.

4.3.4 Oxygen Transfer Experiment

The static method of gassing-out was employed to determine the kra values of xanthan solutions in various reactor systems. The estimation of kLa by this technique depended upon monitoring the increase in dissolved oxygen concentration of a solution during aeration. According to Van't Riet (1979), the use of commercial available electrodes with a response time of 2 to 3 seconds should enable a kLa of up to 360 h_1 or 0.1 s'1 to be measured accurately (Stanburg and Whitaker, 1984). The dissolved oxygen in the solution was first removed by flushing the liquid with nitrogen gas. The deoxygenated liquid was then aerated, and the increase in the dissolved oxygen concentration was monitored with a 121 dissolved oxygen (DO) probe connected to a DO meter (New Brunswick DO-40, New

Brunswick Scientific Co., Edison, NJ), which was connected to a chart recorder

(LINSEIS L-4000 Digital Flatbed Recorder, Electronic Development Co., Boston, MA) to record the time course of dissolved oxygen tension (DOT) in the broth. The experiment was repeated for each condition studied.

4.3.5 Determination of C* and Cl

The dissolved oxygen concentrations in the broth were calculated from DOT values.

The dissolved oxygen concentration in solution can be assumed to be proportional to DOT, with 100%-DOT in water equivalent to the solubility of oxygen in water. The water solubility of oxygen at 1 atm pure oxygen at 30°C is 1.16 mM/L (Bailey and Ollis, 1986).

Since total pressure is usually not far from atmospheric pressure, proportionality of solubility and partial pressures (Henry's law) can be assumed without introducing appreciable errors. The solubility of oxygen at 1 atm air in water at 30°C can be determined from the volumetric percentage of oxygen in air, i.e. 21%, and should be 0.2436 mM/L.

The solubility (C*) and concentration of dissolved oxygen (Cl) in the xanthan fermentation broth were determined from the measured DOT values (in percentage) times the solubility of oxygen in pure water (0.2436 mM/L).

4.3.6 Determination of kLa

Integration of equation (1) yields:

In (C* - C/,) = 0 - (k,a)t. (4.4) 122 where

4.4 RESULTS AND DISCUSSION

4.4.1 C* in Xanthan Broth

Oxygen solubility was reported to be affected by several parameters, including the composition of the medium {Schumpe et a i, 1982). As shown in Figure 4.2, the DOT

level at the saturation state decreased with increasing the xanthan concentration. It is clear that the saturated dissolved oxygen concentration { C * in mmoles dm 3) in the xanthan

broth decreased with increasing the xanthan concentration. The solubilities of oxygen at

various xanthan concentrations determined from the DOT in comparison to that in pure

water are shown in Figure 4.3. As shown, C* decreased rapidly form 0.2436 mM/L in water to 0.185 mM/L in 3.5% xanthan solution. Thus, OTR would also decrease with increasing xanthan concentration even if Jfc/a was maintained at a constant value.

4.4.2 kLa in STR-DT

Figure 4.4 shows the effects of agitation speed on kLa for the various systems studied. In general, k^a increased with increasing the agitation speed. As shown in Figure

4.4(a), in the conventional STR-DT system, the kLa value varied significantly with the agitation speed at low xanthan concentrations, e.g. 12 g/L. However, this effect decreased as the xanthan concentration increased. As the xanthan concentration increased to 35 g/L, the difference in the kyi value at various agitation speeds became insignificant, indicating IUE 4.2: FIGURE

Dissolved Oxygen Tension (%) 100 40 20 80 60 0 Typical dissolved oxygen tension profiles obtained with water and and water with obtained profiles tension oxygen dissolved Typical xanthan broth in the static gassing-out experiment gassing-out static the in broth xanthan 100 ae 12 Xnhn .% Xanthan 2.2% Xanthan 1.2% Water Air 200 Nitrogen 0 700 300 ie (s) Time 0 0 800 600 400 500 TT m IUE 4.3: FIGURE

C* (mM/L) feto ata cnetaino xgnslblt C) n xanthan in (C*) solubility oxygen on concentration xanthan of Effect 0.18 0.20 0.19 0.22 0.23 0.21 0.25 0.24 broth at 30 at broth 0 Xanthan Concentration Concentration Xanthan °C. 10 20 30 (g/L) 40 124 FIGURE 4.4: FIGURE

m a (1/s) k a (1/s) 3 0.02 0.00 0.01 0.04 0.06 0.03 0.05 0.00 0.02 0.01 0 0 ■ STR-DT STR-MP kLa in various systems: (a) STR-DT; (b) STR-MP; (c) STR-WIO; and (d) (d) CPBR-LC. and STR-WIO; (c) STR-MP; (b) STR-DT; (a) systems: kLavarious in 0 40 0 80 1000 800 600 (rpm) Speed 400 Agitation 200 0 40 0 60 1000 600 600 400 200 (b) (a) Water; 2.4% 2 . 2 % 0.06 0.03 0.05 0.04 0.00 0.02 0.01 0 STR-WIO CPBBR-LC I ' Water 0 40 0 80 1000 800 600 400 200 oain pe (rpm) Speed Rotation 0 20 0 400 300 200 100 (d) (C) ■*w/o. Water 3.5% 2 3.5% 1 1 . . . 2 2 2 % . % ■ % 500 126 that increasing agitation was no longer effective in increasing oxygen transfer al this high xanthan concentration.

4.4.3 kLa in STR-MP

The axial flow pattern in STR-MP system resulted in the formation of larger air hubbies, which deteriorated oxygen transfer in viscous xanthan solutions. As shown in

Figure 4.4, k^a in STR-MP was approximately Five to six times lower than that in STR-

DT. Poor mixing in the wall region at high xanthan concentrations might have also resulted in local oxygen limitation, and this presented another serious problem in STR-MP.

However, k^a did not seem to be greatly affected by the xanthan concentration in the STR-

MP.

4.4.4 kLa in STR-WIO

Ju and Zhao (1993) reported that at oil fractions exceeding 10 vol%, STR-WIO shows less pseudoplastic flow behavior and possesses considerably lower effective viscosity. According to our experience, the water-in-oil dispersion could be maintained with up to 60% of aqueous phase present in the emulsion without phase inversion. At a xanthan concentration of 35 g/L in 50% water-in-oil system, the diameters of the droplets formed were observed to be in the range of only 0.01 mm at 800 rpm to 0.05 mm at 4(H) rpm. Thus, the liquid-liquid interfacial area was high at 800 rpm and a dramatic increase in kia was obtained in STR-WIO as compared with STR-DT (Figure 4.4(c)).

However, in spite of the high kia values and good mixing characteristics exhibited in STR-WIO, agglomeration of the xanthan droplets was observed. This observation was 127 in agreement with that reported by Schumpe e ta i (1991). The agglomeration decreased the product yield similar to the effect of dead zones caused by yield stress in a normal fermentation. A more serious problem of emulsion fermentation is the contamination of the product by the emulsifier. This could be overcome by complicated separation procedures that could hardly be avoided during production. For the 3.5% xanthan-in-oil emulsion, at least 6 hours at a standstill were required to separate the two phases. As suggested by

Schumpe et a i (1991), a crude product might be obtained by direct spray-drying of a suitable emulsion. However, the difficulties associated with product recovery by emulsion breaking and removal of the emulsifier are crucial with respect to the selection of suitable oil and emulsifier.

It is also noted that the kia value in STR-WIO may be higher than that in STR-DT, the actual medium volume was only 50% of the total phase volume in the reactor. Thus, the productivity based on the total reactor volume could be lower even though OTR in the medium was enhanced.

4.4.5 kLa in CPBR-LC

For CPBR operated in the liquid-continuous mode with a recirculation stream, the kia value increased with increasing the rotational speed for all xanthan concentrations studied (Figure 4.4(d)). It was observed that the centrifugal container with tilted baffles was able to interact with vessel baffles to create turbulent mixing, which subsequently provided the bioreactor with small air bubbles. The higher the rotational speed, the more turbulent the flow and the better the oxygen transfer. The interfacial area, a , increased with increasing the rotational speed, giving higher kia values at higher speeds. However, at agitation speeds higher than 350 rpm, there was no further increase in the kia value for 128 xanthan solutions with concentrations lower than 2.2%. At the rotational speed higher than

350 rpm, severe agitation and vibration resulted in the formation of large bubbles and

foaming. The kia value at 3.5% xanthan solution could be higher at a rotational speed

higher than 500 rpm. However, for the present system 500 rpm was the maximum speed

advisable for safe operation "f the reactor. Nevertheless, at the same rotational speed, kja

values for CPBR were much higher than those for other bioreactor systems.

4.4.6 kLa in CPBR-GC

It is desirable to know how efficient the rotating fibrous matrix can be used as a

gas-liquid contacting device for oxygenating the viscous xanthan broth. The pseudoplastic

(shear-thinning) property of xanthan solutions allows efficient pumping of xanthan

polymers at high pumping (shear) rates. It is known that at a high rotational speed or under

high gravitational force (> 2-3 g) conditions, gas and liquid can be passed through a

contacting medium at high rates without flooding to achieve a high mass transfer rate that is

2 or 3 orders of magnitude higher than that obtained in conventional mass transfer equipment (Mohr and Khan, 1987; Fowler and Balasundaram, 1989; Smelser et a i, 1990).

In this study, however, the liquid recirculation rate was limited to 270 mL/min because of

the available pump capacity.

Since air was provided from the bottom of the vessel, the system was aerated

through both the 2.5 liters broth in the reactor vessel with agitation and the recirculating stream in contact with air in the rotating fibrous matrix. In order to evaluate the aeration efficiency of the rotating fibrous matrix, the CPBR was operated either with or without recirculation to determine the corresponding kia values. The difference between the kia values obtained from these two operating conditions can be attributed to the direct gas- 129 liquid contact in the fibrous matrix. Figure 4.5 shows the results for various xanthan

solutions at various rotational speeds. It was found that the amount of oxygen transferred

due to the recirculation varied with both the xanthan concentration and the rotational speed.

At low xanthan concentrations, the recirculation stream was effective only at low rotational

speeds. This was because proper agitation and air sparging were sufficient to aerate the

liquid solution. On the other hand, for higher concentration xanthan solutions, kia was

much higher with recirculation than without recirculation. This is because at high xanthan

concentrations the ability of aeration by conventional agitation decreased, resulting in more

significant effect from the direct air-liquid contact in the rotational fibrous matrix. Thus,

the rotating fibrous bed in CPBR can be an effective aeration device for viscous xanthan

broth. The kia value should be higher at a higher recirculation rate.

4.4.7 Comparison of Various Bioreactor Systems

Figure 4.6 compares the maximum kia and OTR obtained at the maximum

rotational speed in all bioreactor systems studied. STR-MP was found to be the poorest

system at all xanthan concentrations studied. At xanthan concentrations below 3%, STR-

WIO gave the highest kia values and OTR among all systems studied. However, as aforementioned, practical difficulties exist in recovering xanthan polymers from such an emulsion system. CPBR was found to have lower kia values than those of conventional

STR-DT at low xanthan concentrations (< 2%). However, the problem was not as serious as it appeared to be, because at such low xanthan concentrations there was no oxygen limitation for xanthan fermentation. CPBR had higher kia than STR-DT at xanthan concentrations > 2%, and comparable kia values to that for STR-WIO at 3.5% xanthan concentration. Furthermore, CPBR seemed to be less sensitive to the xanthan concentration than other systems in the kia value, suggesting that CPBR can be used more

f) B B 2.2% (w/) O 2.2% (w/o) B 3.5% (W O 3.5% (w/o) (C) Rotation Speed (rpm) 100 200 300 400 500 600 0 100 200 300 400 500 600 0.02 0.01 0.00 0.04 0.01 0.00 0.03 0.03 0.02 - - — i (w/) (w/o) — ■— » » ■ ■ i Water Water __ S ’ f l ° ■ 1.2% (w/) O 1.2%(w/o) — *— b O O Water i : s * ' (a) (b) <- <- D o — 1 a B ■o — 1— Rotation Speed (rpm) 1 J i I ■ ■ • ■ — 1 ■o ■o ° - - - o ■ II ■ 0 100 200 300 400 500 600 0 100 200 300 400 500 600

in CPBR-GC kLa at various xanthan concentrations: 0%; (a) (b) 1.2%; (c) 2.2%; and (d) 3.5%. 0.01 0.01 0.03 0.00 0.05 0.03 0.04 0.01 0.00 0.02 0.04 0.05

(S/I) B 1 3| (S/I) e J 0.02 X FIGURE4.5: I E 4.6: RE U FIG

OTR (g/L h) Max. k a (1/s) Comparison of maximal k^a and OTR in various bioreactor systems. bioreactor various in OTR and k^a maximal of Comparison 0.03 0.01 0.05, 0.00 0.02 0.06 0.04 0.5 0.0 0 1 2 3 40 30 20 10 0 ata Cnetain (g/L) Concentration Xanthan 10 wo (b) 20 (a) □T 30 40 131 132 efficiently than other systems at higher xanthan concentrations. Similar trends were found for OTR among all systems studied (Figure 4.6(b)). The OTR was estimated using the determined and C* values at the corresponding xanthan concentration and the C7, value assumed at 20% DOT.

With the improved oxygen transfer character in CPBR at high xanthan concentrations, the CPBR system is practically the preferred system for xanthan fermentation in terms of oxygen transfer ability without problems associated with product recovery.

4.5 CONCLUSION AND RECOMMENDATION

Both kjja and OTR decrease with increasing the xanthan concentration. For all systems studied, STR-WIO provided the highest OTR for the xanthan concentration up to

3%. STR-MP was not appropriate for xanthan fermentation due to its poor aeration performance. STR-DT was also poor when the xanthan concentration was 3.5% or higher.

CPBR provided good OTR at all concentrations studied and was the least sensitive system to xanthan concentration. The oxygen transfer rate in CPBR-GC was the highest among various systems studied at 3.5% xanthan concentration. Therefore, the CPBR system is the preferred system for xanthan gum fermentation in terms of oxygen transfer ability. At high xanthan concentrations, the ability of aeration by conventional agitation decreased.

The direct air-liquid contact provided in the rotational fibrous matrix thus becomes an important mechanism for aeration in CPBR. Also, it is more important to use a high rotational speed for high concentration xanthan solutions. A better constructed rotating fibrous matrix holder that can be operated at a higher speed than that has been achieved in this study will be needed to provide better oxygen transfer. Increase in the flow rate in the 133 recirculation stream, which provided direct contacts of gas, liquid, and matrix (cells) in

CPBR. should be able to further enhance the oxygen transfer rate and xanthan productivity in the reactor.

4.6 REFERENCES

Ahmad, M.N.; Holland, C.R.; McKay, G. Mass transfer studies in batch fermentation: mixing characteristics. J. Food Eng. 1994,23, 145-158.

Behrens, U.; Klima, M.; Fiedler, S. Growth and accumulation of polysaccharide by Xanthomonas campestris. Z Allg. Mikrobiol . 1980, 20, 209-213.

Bhavaraju, S.M.; Russel, T.W.F.; Blanch, H.W. The design of gas sparged devices for viscous liquid systems. AlChE J. 1978, 24(3), 454-466.

Cottrell, I.W.; Kang, K.S.; Kovacs, P. Xanthan Gum. In Handbook of Water-Solubfe Gums and Resins: ; Davidson, R. L. (Ed.), 1980; McGraw-Hill: New York; Chapter 24.

Flores, F.; Torres, L.G.; Galindo, E. Effect of the dissolved oxygen tension during cultivation of X. campestris on the production and quality of xanthan gum. J. Botechnoi 1994, 34, 165-173.

Funahashi, H.; Maehara, M.; Taguchi, H.; Yoshida, T. Effects of agitation by flal-bladed turbine impeller on microbial production of xanthan gum. Journal of Chemical Engineering o f Japan, 1987,20(7), 16-22.

Galindo, E.; Nienow, A.W. Mixing of highly viscous simulated xanthan fermentation broths with the Lightnin A-315 impeller. Biotechnol. Prog . 1992, 8, 233-239.

Gonzales, R.; Johns, M. R.; Greenfield, P. F.; Pace, G. W. Xanthan gum precipitation using ethanol. Process Biochemistry, 1989, 24(6), 200-203.

Henzler, H.J. Engineering design data for stirred vessels as fermentors. Chem. Ing. Tech. 1982, 54(5), 461-476.

Herbst, H.; Peters, H.-U.; Suh, I.-S.; Schumpe, A.; Deckwer, W.-D. Oxygen transfer during xanthan fermentation. 1987 AIChE Annual Meeting, 175a; New York, 1987.

Herbst, H.; Schumpe, A.; Deckwer, W.-D. Xanthan production in stirred tank fermentors: oxygen transfer and scale-up. Chem. Eng. Technol. 1992, 75, 425-434.

Herbst, H.; Suh, I. S.; Peters, H. U.; Schumpe, A.; Deckwer, W. D. Comparison of various fermentor types used for production of xanthan. DECHEMA Biotechnol. Conference. 1989, 3, 495-498. 134 Holzwarth, G. Molecular weight of xanthan polysaccharide. Carbohyd. Res. 1978,66, 173-186.

Jeanes, A. Applications of extracellular microbial polysaccharide-polyelectrolytes: Review of literature, including patents. J. Polym. Sci. Symp. 1974,45, 209-227.

Ju, L.-K.; Ho, C.S. Oxygen diffusion coefficient and solubility in n-hexadecane. Biotechnol. Bioeng. 1989,34. 1221-1224.

Ju, L.-K.; Zhao, S. Xanthan fermentations in water/oil dispersions. Biotechnol. Techniques 1993,7(7), 463-468.

Kennedy, J. F.; Bradshaw, I. J. Production, properties and applications of xanthan. In Progress in Industrial Microbiology, vol. 19. Bushell, M.E. (Ed.), 1984; Elsevier Science, New York, NY; pp. 319-371.

Lilly V.G.; Wilson, H.A.; Leach, J.G. Bacterial polysaccharides: (II) Lab.-scale production of polysaccharides by species of Xamhomonas. Appt. Microbiol. 1958, 6, 105-108.

Lim, T.; Uhl, J.T.; Prud’homme, R.K. Rheology of self-associated concentrated xanthan solutions. J. Rheology , 1984, 28(4 J, 367-379.

Lo, Y.M.; Yang, S.T.; Min, D.B. Xanthan gum fermentation by immobilized Xanthomonas campestris in a novel packed-bed, centrifugal bioreactor. 1995, Unpublished.

Margaritis, A.; Pace, G.W. Microbial Polysaccharides. In Comprehensive Biotechnology, vol. 3; Moo-Young, M. (Ed.), 1985; Pergamon Press, New York, NY; pp 1005-1043.

Margaritis, A.; Zajic, J.E. Mixing, mass transfer, and scale-up of polysaccharide fermentations. Biotechnol. Bioeng. 1978,20(7), 939-1001.

McNeely, W.H.; Kang, K.S. In Industrial Gums. Whistler, R.L. (Ed.), 1973; Academic Press, New York, NY; pp 476-497.

Misra, T.K.; Barnett, S.M. Evaluation of a novel foam fermentor in the production of xanthan gum. Biotechnol. Progresses, 1987,227-237.

Moraine, R. A.; Rogovin, P. Kinetics of the xanthan fermentation. Biotechnol. Bioeng. 1973, 15, 225-237.

Mukhopadhyay, S.N.; Das, D.K. In Oxygen Responses, Reactivities, and Measurements in Biosystems. 1994; CRC Press, Ann Arbor, MI; pp 127-178.

Nakajima, S.; Funahashi, H.; Yoshida, T. Xanthan gum production in a fermentor with twin impellers. J. Ferment. Bioeng. 1990,70, 392-397.

Pace, G.W.; Righelato, R.C. Production of extracellular microbial polysaccharides. In Advances in Biochemical Engineering, vol. 15; Fiechter, A. (Ed.), 1980; Springer- Verlag; New York; pp 41-70. 135 Peters, H.-U.; Herbst, H.; Hesselink, P.G.M.; Lunsdorf, H.; Schumpe, A.; Deckwer, W.-D. The influence of agitation rate on xanthan production by Xanthomonas campestris . Biotechnol. Bioeng. 1989,34, 1393-1397.

Peters, H.-U.; Suh, I.-S.; Schumpe, A.; Deckwer, W.-D. Modeling of batchwise xanthan production. The Canadian Journal of Chemical Engineering, 1992, 70 , 742-750.

Pinches, A.; Pattern, L.J. Rate and yield relationships in the production of xanthan gum by batch fermentation using complex and chemically defined media. Biotechnol. Bioeng. 1986 ,2 8 , 1484-1496.

Schumpe, A.; Quicker, G.; Deckwer, W.D. Gas solubilities in microbial culture media. Adv. Biochem. Eng. 1982, 24, 1-38.

Schumpe, A.; Diedrichs, S.; Hesselink, P.G.M.; Nene, S.; Deckwer, W.D. Xanthan production in emulsions. Biochemical Engineering - Stuttgart, (Proceedings. International Symposium on), 2nd, Meeting Date 1990, 196-9, 1991.

Shu, C.H.; Yang, S.T. Effect of temperature on cell growth and xanthan production in batch cultures of Xanthomonas campestris. Biotechnol. Bioeng. 1990, 35, 454-468.

Sittig, W. Fermentation reactors. CHEM TECH, 1983, 13(10), 606-613.

Solomon, J.; Elson, T.P.; Nienow, A.W. Cavern sizes in agitated fluids with a yield stress. Chem. Eng. Commun. 1981, 11, 143-164.

Stanbury, P.F.; Whitaker, A. Aeration and agitation. In Principles of Fermentation Technology. Pregamon Press: New York; 1984, Chapter 9, pp 169-192.

Suh, I.S.; Schumpe, A.; Deckwer, W.D. Xanthan production in bubble column and air-lift reactors. Biotechnol. Bioeng. 1991, 39, 85-94.

Van't Riet, K. Review of measuring methods and results in non-viscous gas-liquid mass transfer in stirred vessels. Ind. Eng. Chem. Process Des. Dev. 1979, 18(3), 357-360.

Weiss, R.M.; Ollis, D.F. Extracellular microbial polysaccharides. I. Substrate, biomass, and product kinetic equations for batch xanthan gum fermentation. Biotechnol. Bioeng. 1980, 22, 859-873.

Whitcomb, P.J.; Macosko, C.W. Rheology of xanthan gum. J. Rheol. 1978, 22(5), 493- 505.

Yang, S.T.; Zhu, H.; Li, Y. Continuous propionate production from whey permeate using a novel fibrous bed bioreactor. Biotechnol. Bioeng. 1994, 43(11), 1124-1130.

Zaidi, A.; Ghosh, P.; Schumpe, A.; Deckwer, W.-D. Xanthan production in a plunging jet reactor. Appl. Microbiol. Biotechnol. 1991, 35, 330-333. CHAPTER V

ULTRAFILTRATION OF XANTHAN GUM FERMENTATION BROTH

5 J SU M M A RY

Ultrafiltration of xanthan gum solution as an alternative method to alcohol precipitation for xanthan gum recovery from dilute fermentation broth was studied. A polysulfone membrane hollow fiber (with 500,000 MWCO) tubular cartridge was used.

The xanthan fermentation broth, which is highly viscous at its normal concentration of -2.5

(w/v) %, was concentrated to 13.5 (w/v) % or higher, with a recovery yield of -95% or higher. During ultrafiltration, the flux remained almost constant for xanthan concentrations up to -8%. It then decreased dramatically as the xanthan concentration increased beyond

8%. The decreased filtrate flux can be attributed to the decreased pumping (shear) rate and the increased viscosity that resulted at higher xanthan concentrations. When the xanthan concentration was held constant during ultrafiltration, the filtrate flux remained almost unchanged for the entire two-hour period studied, suggesting that the process was stable and no significant fouling occurred. In general, the filtrate flux decreased with increasing xanthan concentration and increased with increasing pumping (shear) rate and trans­ membrane pressure difference. Changing the solution pH had a slight effect on the viscosity of xanthan solution, but did not affect the filtration performance. Even under

136 137 high-shear-rale conditions, ultrafiltration did not give any adverse effects on the rheological properties and molecular weight of the xanthan polymer. Thus, ultrafiltration can be used to concentrate xanthan broth from fermentation by a factor of four or higher, thereby reducing the amount of alcohol needed for xanthan recovery by at least 80%.

Keywords: ultrafiltration, xanthan gum, polysaccharide, kinetics

5.2 INTRODUCTION

Many industrially important polysaccharides, such as xanthan gum and dextran, are or can be produced from sugars via microbial fermentations (Margaritis and Pace, 1985).

Xanthan gum, with a reported molecular weight ranging from 3-7.5 million (Norton et at .,

1984) to 15-50 million (Dintzis etal., 1970; Holzwarth, 1978), is extensively used in food, agricultural, and oil-recovery industries as a thickening, stabilizing, and suspending agenl

(Cottrell et at., 1980). The total consumption of xanthan gum in the U.S. is estimated at over 20 million lbs per year, with an annual growth rate of 7%.

The present industrial process for xanthan gum production is energy-intensive and costly. This is mainly because the high viscosity of xanthan gum causes the agitation and aeration in the conventional agitated tank fermentor to be extremely difficult and consequently limits the final xanthan concentration from fermentation to below -3 (w/v) %.

The recovery of xanthan gum from fermentation broth is generally done with alcohol precipitation. Two to three volumes of alcohol are added to one volume of the broth. The large amount of alcohol required in the recovery process is costly, even with nearly complete recovery of the alcohol by distillation (Gonzales et at., 1989). Distillation of alcohol not only consumes large amounts of energy, but also causes air pollution. A new, 138 innovative production process that is energy efficient, environmentally benign, and cost

effective is thus needed.

There have been many attempts to increase xanthan concentration and to lower the

energy costs by using new types of bioreactors or agitation designs with improved aeration

and oxygen transfer for viscous xanthan fermentation (Herbst et al., 1989; Suh et al.,

1991; Nakajima et al., 1990; Robinson and Wang, 1988; Ju and Zhao, 1993). Some of

these fermentation studies have reported a xanthan concentration of 5% or higher.

However, none of these fermentation improvements was able to economically produce a

xanthan concentration of 10% or higher for industry use. Other recent research efforts thus

have focused on either improving alcohol precipitation (Gonzales et al ., 1989; Flahive et

al., 1994) or developing new post-fermentation concentration and recovery methods, such

as the ultrafiltration process reported in this article.

Ultrafiltration is a well established, low pressure driven, mechanical process for

simultaneously purifying, separating, and concentrating suspended solids, colloids, and

high-molecular-weight materials in solution. In ultrafiltration, liquid is usually pumped

across the membrane surface at high velocity to prevent cake formation, resulting in high

filtration rates that can be maintained continuously. Although ultrafiltration has been

widely studied and used in industry for protein concentration and waste stream clean-up,

very little work has been done with polysaccharide solutions (Barker and Till, 1992; van

Oers et al., 1992; Yeh and Cheng, 1993), and no report can be found in the open literature

for its application to concentrating highly viscous xanthan gum solutions.

The high viscosity of xanthan solution at low concentrations presents a major challenge in using ultrafiltration to concentrate xanthan broth, as the filtrate flux generally decreases rapidly with increasing the solution viscosity (Wang, 1988; Gill et al., 1988).

However, the xanthan solution shows a high degree of pseudoplasticity, i.e., the viscosity 139 decreases rapidly as the shear rate increases (Kang and Pettiu, 1993). This shear-thinning property should allow efficient ultrafiltration of xanthan polymer at high pumping (shear) rates. One objective of this study was thus to determine if the ultrafiltration process could be efficiently used to remove water and impurities in the xanthan fermentation broth to produce a high concentration of purified xanthan solution.

In membrane filtration, the volumetric flux of permeate, Jv, usually can be modeled by the following equation (Fane and Radovich, 1990):

A P - A n AP J - (5.1)

where AP is the transmembrane pressure drop, A n is the osmotic pressure difference across the membrane, rj is the solution viscosity, and ^ / f ( is the total resistance of filtration. In ultrafiltration of macromolecules such as xanthan polymer, An is negligible as compared to AP. The solution viscosity is dependent on the concentration of xanthan gum and the shear rate.

In this work, we studied the feasibility of using a hollow fiber ultrafiltration unit for concentrating dilute xanthan broth produced from fermentation. Several process parameters, including xanthan concentration, transmembrane pressure difference, pumping

(shear) rate, and solution pH, were studied to evaluate their effects on filtration performance. Also, possible shearing effects from ultrafiltration on the xanthan polymer were investigated by examining the rheological properties and molecular weight of xanthan gum before and after ultrafiltration. 140

5.3 EXPERIMENTS

5.3.1 Xanthan Broth

Xanthan broth was produced from fermentation of a glucose-containing medium using Xanthomonas campestris NRRL B-1459. The medium composition and fermentation procedures have been described elsewhere (Shu and Yang, 1990). Unless otherwise noted, the xanthan broth from fermentation was heat-treated at 60°C for -30 minutes to lyse cells before use in this study. A commercial xanthan gum (TIC Gums,

Inc., Belcamp, MD) was also used in this study for comparison purpose. The pH value of the xanthan solution was ~7, unless otherwise specified.

5.3.2 Ultrafiltration System

A tubular polysulfone membrane hollow fiber ultrafiltration cartridge (Model No.

0720199, Koch Membrane Systems Inc., Ann Arbor, MI) was used in this study. The ultrafilter contained 18 hollow fibers, each has 40 cm long and 106 mil or 2.69 mm diameter, and had a total membrane area of 0.7 ft2. According to the manufacturer, the molecular weight cut off (MWCO) of the membrane was 500,000.

Figure 5.1 shows a schematic diagram of the ultrafiltration system. A peristaltic pump (Millipore XX-80-000-00) was used to pump the broth through the tube (lumen) side of the hollow fiber cartridge at a selected high speed. The tubing used was 0.24 in. (6.4 mm) I.D. tubing (Masterflex 6402-24). The actual pumping (flow) rate ranged from 100 mUmin to 1000 mL/min, depending on the xanthan concentration, pressure, and pump speed used. The pressure at the outlet of the retentate stream was monitored with a pressure gauge and was regulated with a valve. This pressure was controlled at 25 psig 141

Retentate

Permeate Ultrafiltration hollow fiber module

Pump Xanthan gum solution

FIGURE 5.1: Hollow-Fiber Ultrafiltration system used in this study. 142 unless otherwise stated. The highest operating pressure recommended by the manufacturer for this ultrafilter is 30 psig. The system was operated in batch mode with the retentate returned to the feed tank, which was a 5-liter fermentor (Marubishi MD-300) maintained at a constant temperature of 30°C. In the feasibility study, the permeate was collected in a separate tank or in a graduated cylinder to monitor its flow rate. In the kinetic studies of various process parameters, the permeate stream was recycled back to the feed tank to keep a constant xanthan concentration in the feed. During the experiment, the permeate flow rate was measured by collecting the permeate in a graduated cylinder for a timed period.

After each batch ultrafiltration, the hollow fiber membrane cartridge was cleaned by recirculating a cleaning solution containing sodium hypochlorite (prepared from 200 ppm chlorine solution by adding NaOH to pH 10.5-11). The water flux was then checked to ensure the membrane had been properly cleaned to restore its original performance,

5.3.3 Process Feasibility Study

Batch ultrafiltration of xanthan broth with initial concentrations up to 4.5 (w/v) % was studied. The initial total liquid volume was 3 liters. The xanthan broth was concentrated until no water in the xanthan broth could be efficiently removed as permeate.

At proper time intervals, the collected permeate volume was measured and used to estimate filtrate flux and the xanthan concentration in the retentate. The filtrate flux, which is the membrane capacity per unit area and unit time, was determined from the measured permeate flow rate divided by the total available hollow fiber membrane area. The initial and final concentrations of xanthan gum in the broth were determined by alcohol precipitation followed by dry weight measurement. The viscosity of the permeate was also measured to 143 estimate the amount of xanthan gum present in the permeate. A material balance was then done to determine the recovery yield of xanthan gum from ultrafiltration.

5.3.4 Kinetic Studies

The effects of various process parameters, including xanthan concentration, shear or liquid flow rate, transmembrane pressure, and the solution pH, on ultrafiltration performance were studied with the concentration of xanthan broth being held at constant by recycling both the retentate and permeate streams back to the feed tank.

5.3.4.1 Membrane fouling

Membrane fouling phenomena were first studied at three different xanthan concentrations (2.7%, 6.6%, and 10.0% wt/v) for a period of 2 hours. The permeate flow rate was monitored throughout each run. Membrane fouling was then studied with xanthan fermentation broth without any heat treatment. One study used the regular fermentation broth that contained cells, and a second study used cell-free fermentation broth produced from an immobilized cell fermentation (Chapter III). The fouled membrane surfaces were examined by scanning electron microscopy.

5.3.4.2 Effects of transmemhrane pressure

The effect of trans-membrane pressure difference on filtrate flux was studied with water and 2.5% xanthan solution at three different pressures (appr. 10, 20, and 27 psig).

5.3.4.3 Effects of xanthan concentration

A wide range of xanthan concentration from 0.03% to 10% was studied to evaluate the concentration effect on filtrate flux at a fixed pumping rate of -1000 mL/min, which 144 gave a shear rate of -500 s 1 on the surface of the hollow fiber membrane. The apparent

viscosity of the xanthan solution at this high shear rate was measured using a Rheometrics

Asphalt Analyzer (RAA, Rheometrics Inc., Piscataway, NJ).

5.3.4.4 Effects of pumping or shear rate

The pumping rate was varied from 100 mL/min to 1000 mL/min to study the effect

of shear rate on ultrafiltration performance. The transmembrane pressure was fixed at 27 psig and xanthan concentration at 2.7%. The shear rate (f) at the wall inside the hollow

fiber tubing was calculated from the liquid flow rate in each tubing ( Qj) and the tube radius

(r) using the following equation for laminar flow condition:

r = ^ - « L (5.2, nr mnr

where Q is the total liquid flow rate and m is the number of hollow fibers contained in the

ultrafilter.

5.3.4.5 Effects of pH

The effects of pH were studied between pH 4 and pH 10. The pH value of the

xanthan solution was adjusted by adding either HC1 or NaOH to the desired pH value.

5.3.5 Effects of Shear on Xanthan Polymer

The effects of shear from ultrafiltration on xanthan gum properties (rheological

properties and molecular weight) were also studied. Rheological properties of xanthan solutions were evaluated using a Brookfield viscometer (RVTD II) at various shear rates

(rpm). The molecular weight of xanthan gum before and after ultrafiltration process was 145 determined from the intrinsic viscosity, [rtf, using the Manelkem-Flory-Scheraga equation

(Holzwarth, 1978; Whitcomb and Macosko, 1978).

5.3.6 Analytical Methods

5.3.6.1 Determination of xanthan concentration

The concentration of xanthan gum in the solution was determined by measuring the

solution viscosity and the dry weight from alcohol precipitation (Lo, 1993). Cells in

solution were first removed by centrifugation and then two volumes of ethanol were added

to each volume of the cell-free broth to precipitate xanthan gum. The precipitates were

collected by centrifugation and then dried overnight to determine the dry weight, which was

then used to determine the xanthan concentration in the solution. For xanthan concentrations lower than 0.8% (wt/v), the measured solution viscosity was found to be

proportional to the xanthan concentration. Thus, for most samples, the xanthan concentration was determined from viscosity by comparing it with a correlation line

obtained from standards.

5.3.6.2 Measurement of viscosity

Unless otherwise noted, the apparent viscosity of the xanthan solution was determined using a viscometer (Brookfield RVTD II, spindle #5, at 100 rpm or 6 s-' shear rate). For measurements at a wide range of shear rate (up to 1000 s ’) and xanthan concentration (up to 10%), a Rheometrics Asphalt Analyzer (RAA, Rheometrics Inc.,

Piscataway, NJ) was used. All viscosity measurements were done at room temperature. 146 5.3.7 Scanning Electron Microscopy

After the experiments were done, the hollow fibers were cut open and the membrane surfaces were rinsed with water to remove any excess gum. These samples membrane were fixed in 2.5% glutaraldehyde solution overnight, then were rinsed thoroughly with distilled water and dehydrated gradually with ethanol baths starting at 20% ethanol and proceeding to 100% ethanol. Further drying was performed with liquid C 02.

The dried samples were sputter-coated with 60% gold/40% palladium. SEM photographs were taken using a JOEL Model JSM-820 scanning electron microscope.

5.4. RESULTS AND DISCUSSION

5.4.1 Feasibility Study

Figure 5.2 shows the performance of three ultrafiltration runs with different initial xanthan concentrations. In general, the permeate flux remained nearly constant for a relatively long period until the xanthan concentration reached -6% in the broth. There were some variations in Filtrate flux because the transmembrane pressure drop might not be maintained at a constant level throughout each run. At concentrations higher than 6~8%>, the flux declined rapidly with increasing xanthan concentration. A final concentration of

-13% was reached. The process was stopped because of the difficulty of agitation in the feed tank when the xanthan broth inside was highly concentrated and the volume of the broth was too small to be pumped efficiently.

5.4.1.1 Recovery yield

Based on the initial and final liquid volumes and xanthan concentrations of the retentate in the feed tank, about 95% of xanthan gum was recovered in the retentate from 147

5

4 Xanthan

3

2 Flux

1

0 0 0 50 100 150 200 250 300 350 400 450 (a) 5

4 Xanthan

3

2 Flux 1

0 0 5 7550 100 125 150 175 200 225 Xanthan Xanthan Concentration (%) (b) 5

4 Xanthan

3

2 Flux 1

0 —I o 0 25 50 75 100 125 150 175 200 Time (mln)

FIGURE 5.2: Unsteady-state performance of ultrafiltration of xanthan broth at 25 psig and a fixed pumping rate of 3.5 (-350 mL/min). 148 the ultrafiltration process. The other 5% of xanthan gum was mainly adsorbed or entrapped in the membrane pores. The possible loss of xanthan gum in the permeate was also measured, and was found to be insignificant. The viscosity of the permeate was only

8-10 cP at 6 s-1 shear rate, which was equivalent to a xanthan concentration of less than

0.01% (wt/v). Compared to the initial xanthan concentration of 1% or higher, less than

1 %, if any, xanthan gum could be present in the permeate.

5.4.1.2 Concentration polarization

In ultrafiltration, the filtrate flux was one order of magnitude lower for the xanthan broth than for pure water, possibly because of severe concentration polarization caused by the xanthan polymer. When the xanthan concentration was held constant during ultrafiltration, the filtrate flux decreased slightly with time for the two-hour period studied

(Figure 5.3), indicating that the ultrafiltration process is relatively stable and no significant fouling occurred. This stable performance is not found with protein ultrafiltration, in which fouling and notable flux decline usually continue with time due to continuing protein adsorption and cake formation on the membrane, van Ocrs et al. (1992) studied ultrafiltration of dextran, which is also a polysaccharide but with a much smaller molecular weight than that of xanthan gum, and found that only a polarization layer and no gel layer was built up during filtration. In their experiments, the time to reach steady-state flux for dextran ultrafiltration was less than a minute, whereas it took more than one hour when there was gel layer formation during ultrafiltration of silica particles. Apparently, the adsorption or deposition of xanthan molecules on the membrane and gel layer formation either occurred instantaneously upon contact or was not significant The dramatic decrease in permeate rate at high xanthan concentrations was thus attributed to the concentration polarization effect. I E 5.3: RE U FIG

Jv (10*6m3m-2s*1) taysaepromneo lrflrto fxnhnboh t2 psig 25 at broths xanthan of ultrafiltration of performance Steady-state o . - 0.5 . - 2.5 2 . i ■ i 3.0 1-5 ' 6.55% Xanthan 6.55% ' 1-5 «„«[ .0 - 2.72% Xanthan 2.72% - .0 0 I ■ I and -350 mL/min pumping rate. pumping mL/min -350 and 2 4 6 8 10 120 100 80 60 40 20 0 11 i i * * * * * * ii p h- " i ---- • ■ .... — • 1 __ ------1 * ■ - “ • ------— —. — j 1 Xanthan Xanthan 10.03% 10.03% 1 I 1 l i | * 1 • ■ 1 6.55% Xanthan 6.55% 2,72% Xanthan 2,72% ------ie (min) Time * -----

O 1 1 1 ----- ■ ---- - ■ - . ■---- O ' — ---- - T - - • 1 ----- * ---- b B • ------* It * * —o— -*— i B o 1 * ------—• ■ ■■ j B • ------9> a 4 4 II 149 150 5.4.1.3 Membrane Fouling

The membrane could be fouled by cells and cell debris present in the broth. As shown in Figure 5.4, when the whole fermentation broth containing cells was used for ultrafiltration without prior heat treatment, severe membrane fouling occurred. On the other hand, when the cell-free fermentation broth produced from an immobilized cell fermentation process (Chapter III) was used, the filtrate flux remained almost constant during the entire period studied. It is clear that heat treatment greatly improves the ultrafiltration performance by lysing the cells, but cell debris still might cause some fouling problem over long-term operation. However, the fouled membrane could be cleaned and restored to the original performance by washing with water or a cleaning solution, suggesting that the binding of the xanthan and other solute molecules with the membrane was not loose.

The scanning electron micrographs (Figure 5.5) show the membrane surface fouled with cells from ultrafiltration with the whole xanthan fermentation broth. The surfaces became clean again after washing. In contrast, the membrane surface after ultrafiltration of the cell-free broth was as clean as the unused membrane surface.

5.4.2 Kinetic Studies

The ultrafiltration process must be optimized for several parameters, including xanthan concentration, solution viscosity, pH, pumping (shear) rate, and trans-membrane pressure. As already shown in equation (5.1), the filtrate flux, Jv, should increase with increasing the transmembrane pressure, AP, and decreasing the xanthan concentration and the solution viscosity, /t. Since xanthan gum is a shear-thinning polymer, the viscosity of 151

2.0

1.9 Cell-free 0 —

1.8 Heal treated

1.7

.6

1.5 0 20 4 0 6 0 80 1 0 0 1 20 Time (mln)

FIGURE 5.4; Effect of cells on membrane fouling 152

Original Membrane Surface UF with Whole Broth

UF with Cell-Free Broth Surface After Cleaning

FIGURE 5.5; SEM photographs of membrane before and after use in ultrafiltration of xanthan broth. 153 the xanthan solution decreases with increasing shear rate or liquid velocity, and should result in an increase in the filtrate flux. Since the operating conditions affect the total resistance of the filtration as well as the solution viscosity, it is desirable to combine the viscosity term with the resistance term. Equation (5.1) thus can be reduced to:

yv = — = — ^ (5.3)

where Rlotai >s the total resistance and can be expressed as the sum of the intrinsic membrane resistance, Rm, and the film resistance, Rp which may consist of resistances form gel formation, pore blocking, and concentration polarization. The concentration of xanthan, pH, pressure, and pumping (shear) rate were studied under steady-stale conditions and their effects on filtrate flux and the film resistance are discussed in the following sections.

5.4.2.1 Effect of pressure

The transmembrane pressure difference can be approximated as follows:

AP= ' ~ P p (5.4) 2 where P, and P0 are the inlet and outlet pressure of the retentate stream, respectively, and

Pp is the pressure at the permeate side, which is equal to the ambient pressure. The pressure drop {P, ■ PQ) for flow in the hollow fibers in this system was ~5 psi at all flow conditions studied.

The pressure effect was first studied with pure water. As shown in Figure 5.6(a), the filtrate flux increased with increasing the transmembrane pressure. The intrinsic resistance of membrane, Rm, at various pressures was calculated and is also shown in ] 54

80 Water

70 €.m s ( "M ^01) iu*d I i u c. s (

50

40

30 Rm 20

0.50 0.75 1.00 1.25 1.50 1.75 2.00 zm j 1 ) j #1 q sj

2.5% xanthan Q= 1011.60 mL/min ' * * ■ 1 ■ * ...... 0.50 0.75 1.00 1.25 1.50 1.75 2.0U A P (105 Pa) (b)

FIGURE 5.6: Effects of pressure on filtrate flux and resistance, (a) water (b) 2.5% xanthan solution. 155 Figure 5.6(a). The Rm value was found to be almost unaffected by the pressure and was

-2.4 x 10 9 Pa m2 s n r3.

The effect of pressure on the filtrate flux of 2.5% xanthan solution is shown in

Figure 5.6(b). Again, higher pressure gave a greater flux, but the increase in flux was less

than proportional to the pressure increase because of the increased film resistance at higher

pressures. In general, the higher the transmembrane pressure difference, the more xanthan

molecules buildup at the solution-membrane interface, resulting in a larger film resistance to

permeate flow. During ultrafiltration, once a large molecule has been transported to the

membrane surface the chance of it returning to the bulk solution by diffusion is substantially reduced, especially at high pressures (Jonsson, 1993). It is noted that /fy-was

one order of magnitude larger than Rm.

5.4.2.2 Effect of xanthan concentration

As shown in Figure 5.7, the viscosity of a xanthan solution increased with

increasing the xanthan concentration. The increase in viscosity with concentration was linear when the concentration was lower than 0.6 g/L and was exponential when the concentration was higher than 0.6 g/L.

Figure 5.8 shows the effects of xanthan concentration on the pumping (flow) rate, filtrate rate, filtrate flux, and film resistance at 25 psig and a fixed pump speed. In general, the filtrate flux decreased with increasing the xanthan concentration. The decrease in filtrate rate with increasing xanthan concentration was most profound when the xanthan concentration was lower than 0.1% (Figure 5.8(b)), For concentrations below 0.6 g/L. the decrease with increasing xanthan concentration or solution viscosity was linear. As shown in the logarithmic plot (Figure 5.8(c)), for concentrations between 0.03% and 10%, the 156

1000 Shear Rate = 500 s

9

w 100 9 vt9

10 0 2 3 4 5 6 Xanthan Concentration (%) (a)

20 Shear Rate = 6 s a- 9 15

V) 10 9O VI

5

0 0.00 0.02 0.0 4 0 .0 6 o.oe Xanthan Concentration (%) (b)

FIGURE 5.7: Effect of xanthan concentration on solution viscosity measured at (a) 500 s 1; and (b) 6 s 1 shear rate. I E 5.8: RE U FIG

J, Flltrate Rate (mL/mill) Pump Efficiency (%) fet o xnhn ocnrto o () oa vlmti fo rate flow volumetric total (a) on concentration xanthan of Effects and fouling resistance. The transmembrane pressure was 27 psig. 27 was pressure transmembrane The resistance. fouling and through the hollow-fiber ulirafilter; (b) filtrate rate; and (c) filtrate flux flux filtrate (c) and rate; filtrate (b) ulirafilter; hollow-fiber the through 100 100 100 .01 40 60 80 70 10 90 80 .1 1 .01 ■ 6 0 0 ■

o ata Cnetain (%) Concentration Xanthan 2 2 .1 upRaig 3.5 Reading Pump 400 200 300

100 4 4 (c) (b) (a) 1 6

Pump Reading 9.5 Reading Pump A P = 25 psig 25 A = P Pump Reading 9.5 Reading Pump 8 8 10

10 10 Rl 100 1 12

2 100 .01 .1 10 1

Rf (1®,# Pa mJ m ^s) 157 158 decrease in filtrate flux and increase in film resistance were exponential with increasing the xanthan concentration.

The concentration effect can be attributed to the viscosity effect. As found in a theoretical study by Gill et al. (1988), increasing the solution viscosity would increase concentration polarization, and thus would reduce the filtrate rate. Wang (1988) studied the effect of viscosity on ultrafiltration and found that in the low solution viscosity range, an infinitely small increase in viscosity would cause an abrupt decrease of permeate rate.

Further increases in viscosity would cause moderate decrease in flux, and the viscosity effect becomes insignificant at high viscosity.

The actual pumping (flow) rate also decreased with increasing the xanthan concentration, even at the same pump speed. As shown in Figure 5.8(a), the pumping efficiency or the actual volumetric flow rate at a set pump speed decreased by ~10%-25% when the xanthan concentration increased from 0.1% to -10%. The decrease in flow rale with concentration was linear, and the decreasing rate was higher at lower pumping speeds.

The reduced pumping (flow) rate would result in a lower shear rate, a higher apparent solution viscosity, and a higher film resistance, as discussed further below.

5.4.2.3 Effect of pumping (shear) rate

The shear-thinning characteristics of xanthan broth is shown in Figure 5.9. The viscosity decreases exponentially with increasing the shear rate in the range between 0.1 s ' and 1000 s 1. For shear rales smaller than 0.1 s 1 and greater than 1000 s 1, the effect of xanthan concentration on the apparent solution viscosity was not as significant. It is clear that a sufficiently high shear rate should be used in xanthan ultrafiltration because the solution viscosity becomes formidably high at low shear rates. FIGURE 5.9: FIGURE Viscosity (cP) 102 10 101 3 0 1 4 0 1 10 i l ■ i ■ i >■ il i m i ■ ° V 5 .01 1 1 3.6% I ■

.% * ♦ * - a ■ ~ 0.9% Effect of shear rate on the viscosity of xanthan solution. xanthan of viscosity the on rate shear of Effect .... .- .... .

8 •••. • • • % .1 j ......

ha Rt (1/sec) Rate Shear .. j 1

...... 1 i ■ iii■■ 10 j ■ •...... —i

■ ■■ ■■ 100

1000 159 160 As shown in Figure 5.10, the volumetric flow rate in the hollow Fibers and the shear rate at the membrane surface increased linearly with increasing pumping rate (speed).

The Filtrate rate also increased with increasing pumping rate, but the increase was more profound at lower pumping (shear) rates. For flow rates higher than 200 mL/min (which is equivalent to a wall shear rate of -100 s 1)* the Filtrate rate increased exponentially and the

Film resistance decreased exponentially with the decrease in pumping (shear) rate (Figure

5.10(c)).

It is well known that high shear rate improves the filtrate flux during ultrafiltration because gel layer formation, cake deposition, and membrane fouling are reduced (Jonsson,

1993). However, the observed shear effect on filtrate rate in the present study should be attributed mainly to the viscosity effect on concentration polarization since gel layer formation and membrane fouling were not observed.

The maximum shear rate (at the wall of the hollow Fiber tube) at the highest pumping rate (-1000 mlVmin) studied was -500 s l. At this shear rate, the viscosity of the xanthan solution is on the order of 100 cP. Under this flow condition, the Reynolds number for flow in the hollow fiber tubing was less than 10. The flow was thus laminar in the hollow Fibers studied. Creeping flow conditions could exist with lower pumping rates and higher xanthan concentrations. This unfavorable flow condition may explain the extremely low filtrate flux obtained in ultrafiltration of the highly viscous xanthan solution.

5A2.4 Effect of solution pH

The xanthan molecule contains anionic groups in its side chains at the neutral pH

(Margaritis and Pace, 1985). The solution pH thus may affect the electrostatic charges on the xanthan molecule, which in turn may affect concentration polarization and fouling characteristics of xanthan molecules on the membrane. The solution pH may also affect the 161

1200 1200

1000 1000

800 8 0 0 Mi Row Rate 600 6 0 0 i 06 400 4 0 0 at at 200 45 200 Shear Rate W) 0 0 2 4 6 8 10 (a) 10

8

6

4 A p = 27 psig 2 2.7% Xanthan 0 0 2 4 6 8 10 Pump Reading (b)

100 10 0 tn e 10 r*a <2 1 J v

.1 1 0 100 1000 10000 Flow rate (mL/min) (c)

FIGURE 5.10: Effects of the pumping rate (pump speed) on (a) total volumetric flow rate through the ultrafilter; (b) filtrate rate; and (c) filtrate flux and fouling resistance, (with 3.5% xanthan solution at 27 psig) 162 viscosity of the xanthan solution. As shown in Figure 5.11(a), notable viscosity changes with the pH were found for a commercial TIC xanthan solution (2%), while the viscosity change for the 3.5% xanthan solution produced in our own fermentation laboratory was less than 5%. However, for both xanthan solutions, there was no significant difference in the filtrate rate among various pH values studied (Figure 5.11(b)). It is clear that the solution pH did not have any significant effect on filtrate flux nor on fouling during ultrafiltration of xanthan solution.

S.4.2.5 Filtrate flux

Since the film resistance is dependent on pressure {AP), xanthan concentration (C), and shear rate or flow rate ( Q), equation (5.3) is not convenient to use in engineering design. As already shown by the linear relationship found in the logarithmic plot (Figure 5.8(c) and 5.10(c)), the filtrate flux, Jv (m^nrV1), varied with AP (Pa), C (g/L), and Q

(m3/s) following the empirical equation given below;

Jv = 3.51 x 10 "5 AP025 Q° 32 (5.5) where the proportional constant and exponent values were determined from the logarithmic plots of Jv, vs. AP, C, and Q, respectively. It is noted that the exponents associated with

AP, Q, and C all have an absolute value of less than 1, meaning that their effect on Jv is not linear. Equation (5.5) will be useful in predicting filtrate flux under various conditions, and thus, can be easily used in scale-up design. I E 5.11: RE U FIG

Filtrate Rate (mL/min) Viscosity (cP) 2000 3000 1000 feto h H n a slto vsoiy ad b itae ae(t 3.5 (at rate filtrate (b) and viscosity; solution (a) on pH the of Effect 10 pumping rate and 25 psig). 25 and rate pumping 0 I ■4 i 2 r a ■ 'r ® 8 0 * H 10 pH * ’ ‘ “ pH 4 pH “ ‘ - ■ ' | I ■| » | T*" | ■»" ■| | —I | » I' ■ I 2 4 6 0 0 120 100 80 60 40 20 0 4 8 0 12 10 8 6 4 2 ■ ■ ■ ■ ■ i ■ ■ ■ ■ ■ ■ i ■ ■ I ■ r ...... 2.0% TIC xanthan TIC 2.0% 3.5% Lab xanthan Lab 3.5% .% xanthan 3.5% .% xanthan 3.5% i ie (min) Time . pH (b) (a) i . t. i . p 10 pH * “ pH 4 4 pH “ ° pH 7 7 pH pH ° ° ..... m 164 5.4.3 Effects of Ultraflltration on Xanthan Molecule

The possible effect of shear on xanthan molecules during ultraflltration was investigated by studying the rheological properties and the molecular weight of the xanthan molecules before and after ultraflltration.

5.4.3.1 Rheological properties

Xanthan gum solution is pseudoplastic and its rheological characteristics follow the power law (Galindo et a/., 1989);

r = - K y " (5.6) where r is shear stress, y is shear rate, K is consistency index, and n is flow behavior index. Both K and n can be determined from the logarithmic plot of r vs. y. In this study, the measured torque reading from the viscometer, which is proportional to r, and the rotational speed (rpm) of the spindle, which is proportional to y, was used in the plot shown in Figure 5.12. As can be seen in this figure, the difference between the two xanthan broth samples, one before and the other after ultraflltration, was not significant.

The measured n values were 0.22-0.25. These values were within the range of 0.2 to 0.9 reported for xanthan solutions in the literature (Galindo et a/., 1989).

S.4.3.2 Molecular weight

The molecular weight of a macromolecule can be determined from the intrinsic viscosity, 117], of the solution containing the macromolecule, [rjj is defined as;

(5.7) I E 5.12: RE U FIG

Torque Comparison of rheological characteristics of xanthan solutions before before solutions xanthan of characteristics rheological of Comparison and after ultraflltration. after and .1 k i

n-slope=0.24833 Before UF Before • 1 A

a . 10 rpm =lp=.20 ■ n=slope=0.22108 :

.2 xanthan 2.72% ' fe F : UF After ' ! 100

1000 • 165 166 where C is the xanthan concentration (wt/v) and r}sp = (T} ■ rj0)/r}0. 17 and r}0 are the apparent viscosity of the solution containing the solute and the solvent (water) without the solute, respectively. (nJ can be determined from the intercept of the plot of rfsp/C vs. C.

As shown in Figure 5.13, [ 77] for the xanthan polymer was 1.436 x 10 4 mL/g before ultrafiltralion and 1.597 x 10 4 mL/g after ultrafiltration. These values correspond to molecular weight of 19.1 x 10 6 and 22.4 x 106, respectively. The slight increase in the estimated molecular weight of xanthan polymer was not a surprise since small molecules

(M.W. less than 500,000) were removed in ultraflltration. It is clear that the shear rate under the laminar flow condition did not cause any significant physical damage or alteration of the xanthan polymer. It is thus concluded that the xanthan polymer was not significantly or adversely affected by ultraflltration under the conditions studied.

5.4.4 Energy Consumption

The energy requirement in concentrating the xanthan fermentation broth by ultraflltration is mainly attributed to the pumping. The pressure drop for xanthan broth to flow through the hollow fiber tubes was ~5 psig for the conditions studied. The corresponding power consumption and energy requirement per unit volume of water removed from the broth can be estimated from the pressure drop. In general, the energy requirement for ultraflltration should decrease with increasing AP because of increased filtrate rale at higher AP. Therefore, ultraflltration should be performed at the highest possible pressure. For the hollow fiber cartridge used in this study, the maximum operable pressure was only -30 psig. However, for industrial operation using spiral wound membrane unit, the pressure can go up to 70 psig. I E 5.13: RE U FIG

sp/C (mL/g) 20000 25000 30000 15000 10000 5000 Comparison of the intrinsic viscosity of xanthan solutions before and and before solutions xanthan of viscosity intrinsic the of Comparison 000 0 .0 0 — o after ultraflltration. after Before UF Before After UF After ata Cnetain (g/mL) Concentration Xanthan .39+ + .93+x A = 0.914 RA2 = 6.8953e+5x + 1.4359e+4 .96+ + .80+x A = 0.926' RA2 = 1.6830e+6x + 1.5966e+4 0.001 0.002 167 168 in general, a sufficiently high pumping (shear) rate should be used in xanthan ultrafiltration; otherwise, the solution viscosity may become formidably high for the process. However, the energy requirement per unit volume of water removed from ultrafiltration could be higher at higher pumping rate, even though the filtrate rate is also higher. This is because power consumption increases proportionally but filtrate rate increases less proportionally with increasing the pumping rate. Therefore, there is an optima] pumping rate for minimum energy consumption during ultraflltration. The optimal pumping rate for the process, however, should be determined by also taking into account the membrane cost, which decreases with increasing the pumping rate because of the improved filtrate flux and reduced membrane area required for the process.

The energy consumption in ultraflltration of xanthan gum should be much lower than that used in the current alcohol precipitation method, which requires energy-intensive distillation for alcohol reuse in the recovery process. The energy consumption and cost consideration in ultraflltration of xanthan solution will be further discussed in Chapter VI.

5.5. CONCLUSION

Ultraflltration with the hollow fiber cartridge can be used to concentrate dilute xanthan fermentation broth (-2.5%) to a concentration of -13% or higher, with a recovery yield of greater than 95%. The process is stable and does not harm the qualities of the xanthan polymer. The concentrated xanthan broth would require proportionally smaller amounts of alcohol for xanthan gum production. With a fivefold increase in the xanthan concentration, the amounts of alcohol needed in xanthan recovery would be reduced by

80%. Also, water and solutes present in the permeate from ultraflltration may be recycled 169 for use in fermentation. Compared to the present method for xanthan gum recovery, the ultraflltration process should be more energy efficient and environmentally friendly.

5.6. LIST OF SYMBOLS

C concentration, wt/v or g/L

F filtrate (permeate) rate, mL/s

J v volumetric permeate (filtrate) flux, m 3n r 2s_1 K consistency index in the power law m number of hollow fiber tubing in the ultraflltration unit n flow behavior index in the power law

P Power consumption, W

Pi inlet pressure of the tube side, Pa or psig

Po outlet pressure of the tube side, Pa or psig

Pp pressure at the shell side (permeate outlet). Pa or psig AP transmembrane pressure difference as defined in Equation (5.3), Pa or psig

Q pumping rate or volumetric flow rate, mL/min or m3/s

Qi volumetric flow rate in hollow fiber tubing, mL/min or m3/s r radius of hollow fiber tubing, m

Rf film resistance, Pa m 3m2s m intrinsic membrane resistance, Pa m 3m2s

P total total resistance. Pa n r 3m2s

Y shear rate, s -1 r) viscosity of solution, cP rf0 viscosity of solvent, cP 170 rjsp specific viscosity of solution, unitless

[ rj] intrinsic viscosity of solution, mL/g

T shear stress

it osmotic pressure, Pa

5.7 REFERENCES

Barker, P.E. and Till, A. Using multistage techniques to improve diafiltration fractionation efficiency. J. Membrane Sci., 72: 1-11, 1992.

Cottrell, I.W.; Kang, K.S.; Kovacs, P. Xanthan Gum. In Handbook of Water-Soluble Gums and Resins’, Davidson, R.L. (Ed.), 1980; McGraw-Hill, New York, Chapter 24.

Dintzis, F.R.; Babcock, G.E.; Tobin, R. Studies on dilute solutions and dispersions of the polysaccharide from Xanthomonas campestris NRRL B-1459. Carbohyd. Res., 1970, 13, 257-267.

Fane, A.G.; Radovich, J.M. Membrane systems. In Separation Processes in Biotechnology. Asenjo, J.A. (Ed.), 1990; M. Dekker, New York, NY; Ch. 8 .

Flahive III, J.J.; Foufopoulos, A.; Etzel, M.R. Alcohol precipitation of xanthan gum from pure solutions and fermentation broths, Separation Sci. Technol. 1994, 29, 1673-1687.

Galindo, E.; Torrestiana, B.; Cuernavaca; Garcia-Rej 6 n, A. Rheological characterization of xanthan fermentation broths and their reconstituted solutions. Bioprocess Eng. 1989,4, 113-118.

Gill, W.N., Wiley, D.E., Fell, C.J.D., Fane, A.G., Effect of viscosity on concentration polarization in ultrafiltration, AIChE J., 34, 1563-1567.

Gonzales, R„ Johns, M.R., Greenfield, P.F., and Pace, G.W. Xanthan gum precipitation using ethanol. Process Biochemistry, 24(6), 200-203, 1989.

Ju, L.-K.; Zhao, S. Xanthan fermentations in water/oil dispersions, Biotechnol . Techniq. 1993, 7, 463-468.

Jonsson, A.-S. Influence of shear rate on the flux during ultrafiltration of colloidal substances. J. Membrane Sci., 79: 93-99, 1993.

Herbst, H.; Suh, I. S.; Peters, H. U.; Schumpe, A.; Deckwer, W. D. Comparison of various fermentor types used for production of xanthan. DECHEMA Biotechnol. Conference . 1989,3, 495-498.

Holzwarth, G. Molecular weight of xanthan polysaccharide. Carbohyd. Res. 1978,66, 173-186. 171 Kang, K.S.; Pettilt, D.J. Xanthan, gellan, welan, and rhamsan, In Industrial Gums; Polysaccharides and Their Derivatives; Whistler, R.L.; BeMiller, J.N. (Eds.), 1993; Academic Press, San Diego, CA, pp 341-371.

Lo, Y.-M. Effects of medium composition on product formation and cell growth in xanthan gum fermentation. M.S. Thesis, Ohio State University, 1993.

Margaritis, A.; Pace, G.W. Microbial Polysaccharides. In Comprehensive Biotechnology, vol. 3; Moo-Young, M. (Ed.), 1985; Pergamon Press, New York, NY; pp 1005-1043.

Mulder, M. Basic principles of membrane technology, Ch. 6 . Kluwer Academic Publishers, 198-212, 1991.

Nakajima, S.; Funahashi, H.; Yoshida, T. Xanthan gum production in a fermentor with twin impellers. /. Ferment. Bioeng. 1990, 70, 392-397.

Norton, I.T.; Goodall, D.M.; Frangon, S.A.; Morris, E.R.; Rees, D.A. Mechanism and dynamics of conformational ordering in xanthan polysaccharide. J. Mol. Biol. 1984, 175, 371-394.

Robinson, D.K.; Wang, D.I.C. A transport controlled bioreactor for the simultaneous production and concentration of xanthan gum. Biotechnol. Prog. 1988, 4, 231-241.

Shu, C.H.; Yang, S.T. Effect of temperature on cell growth and xanthan production in batch cultures of Xanthomonas campestris. Biotechnol. Bioeng. 1990, 35, 454-468,

Suh, I.S.; Schumpe, A.; Deckwer, W.D. Xanthan production in bubble column and air-lift reactors. Biotechnol. Bioeng. 1991, 39, 85-94.

Tweddle, T.A., Striez, C., Tam, C.M., and Hazlett, J.D. Polysulfone membranes I. Performance comparison of commercially available ultraflltration membranes. Desalination, 8 6 : 27-41, 1992.

van Oers, C.W., Vorstman, M.A.G., Muijselaar, W.G.H.M., and Kerkhof, P.J.A.M. Unsteady-state flux behavior in relation to the presence of a gel layer. Journal of Membrane Science, 73: 231-246, 1992.

Wang, S.-S., Effect of solution viscosity on ultraflltration flux, J. Membrane Sci., 39 (1988) 187-194.

Whitcomb, P.J.; Macosko, C.W. Rheology of xanthan gum. J. Rheol. 1978, 22(5), 493-

Yeh, H.-M. and Cheng, T.-W. Osmotic-pressure model with permeability analysis for ultraflltration in hollow-fiber membrane modules. Separation Technology, vol. 3: 91-98, 1993. CHAPTER VI

ULTRAFILTRATION PROCESS EVALUATION

6.1 SUMMARY

The energy consumption and operating costs for xanthan recovery from dilute

fermentation broth using ultrafiltralion and alcohol precipitation were studied.

Ultraflltration can be used economically to concentrate xanthan fermentation broth from

2.5% to 15% (w/v) xanthan. Engineering design equations for modeling ultrafiltration of xanthan gum are also developed and used in process analysis. The effects of xanthan concentration, transmembrane pressure, and shear and pumping rates on the performance of ultraflltration were also studied. The optimal condition for maximum filtrate flux and thus minimum membrane costs was determined. It is found that up to 80% of the energy used in recovering xanthan gum from fermentation broth can be saved by using ultraflltration.

Keywords: ultrafiltration, xanthan fermentation, energy consumption.

172 173

6.2 INTRODUCTION

Xanthan gum is a microbial polysaccharide, industrially produced by fermentation

of glucose (Margaritis and Pace, 1985), Because of its high viscosities at low

concentrations, xanthan gum is extensively used in food, agricultural, and oil-recovery

industries as a thickening, stabilizing, and suspending agent (Kang and Pettitl, 1993). The

total consumption of xanthan gum in the U.S. is estimated at over 20 million lbs per year,

with an annual growth rate of 7%.

The present industrial process for xanthan gum production is energy-intensive and

costly. This is mainly because the high viscosity of xanthan solution causes the agitation

and aeration in the conventional agitated tank fermentor to be extremely difficult and

consequently limits the final xanthan concentration from fermentation to below -3% (wt/v).

The recovery of xanthan gum from fermentation broth is usually done by precipitation with

isopropanol. Two to three volumes of alcohol are added to one volume of the broth

(Gonzales etal., 1989; Garcia-Ochoa et al„ 1993; Flahive etal ., 1994). Nearly complete

recovery of the alcohol by distillation for process reuse is essential to process economics.

However, alcohol distillation is energy intensive (Swift and Atkins, 1983).

Ultrafiltration of xanthan gum solutions has been proposed as an alternative method

to alcohol precipitation for xanthan gum recovery from dilute fermentation broth (Lo et at.,

1995). The xanthan solution shows a high degree of pseudoplasticity, i.e., the viscosity decreases rapidly as the shear rate increases (Galindo eta l ., 1989). This shear-thinning property allows efficient ultraflltration of xanthan polymer at high pumping (shear) rates.

The objective of this research was to determine the energy consumption and process economics for the ultraflltration process. 174 In this work, ultrafiltration of xanthan fermentation broth as affected by several process parameters was studied first. The optimal condition for maximum filtrate flux and thus minimum membrane costs was evaluated. Energy consumption for pumping the xanthan solution consists of another major operating cost for the ultraflltration process.

Energy is mainly used to provide the shaft work needed to overcome friction loss or pressure drop for liquid flow in tubes. Pressure drop in pumping the viscous xanthan solution in hollow-fiber tubing was thus studied. The rheological characteristics and flow behavior of viscous xanthan solution were also studied. Finally, the operating costs of xanthan ultrafiltration process were evaluated based on its power consumption in pumping and membrane replacement costs. The energy cost for ultrafiltration was also compared with that for the alcohol precipitation process.

6.3 MATERIAL AND METHODS

6.3.1 Xanthan Broth

Xanthan broth was produced from fermentation of a glucose-containing medium using Xanthomonas campestris NRRL B-1459. The medium composition and fermentation procedures have been described elsewhere (Lo, 1993). Unless otherwise noted, the xanthan broth from fermentation was heat-treated at 60°C for -30 minutes to lyse cells before use in this study. The pH value of the xanthan solution was 7 .

6.3.2 Ultrafiltration

A tubular polysulfone membrane hollow fiber ultrafiltration cartridge (Model No.

0720199, Koch Membrane Systems Inc., Ann Arbor, MI) was used to study the filtrate 175 flux under various conditions. Figure 6.1 shows a schematic diagram of the ultrafiluation

system used in this study. The ultrafilter contained 18 hollow fibers, each was 40 cm long

and 106 mil or 2.69 mm diameter, and had a total membrane area of 0.7 ft2. According to

the manufacturer, the molecular weight cut off (MWCO) of the membrane was 500,000. A

peristaltic pump (Millipore XX-80-000-00) was used to pump the broth through the tube

(lumen) side of the hollow fiber cartridge at a selected high speed. The tubing used was

0.24 in. (6.4 mm) I.D. tubing (Masterflex 6402-24). The actual pumping (flow) rate

ranged from 100 mL/min to 1000 mL/min, depending on the xanthan concentration,

pressure, and pump speed used. The pressure at the outlet of the retentate stream was

monitored with a pressure gauge and was regulated with a valve. This pressure was

controlled at 25 psig unless otherwise stated. The system was operated in batch mode with

the retentate returned to the feed tank, which was a 5-liter fermentor (Manibishi MD-300)

maintained at a constant temperature of 30°C. The permeate stream was recycled back to

the feed tank to keep a constant xanthan concentration in the feed. During the experiment,

the permeate flow rate was measured by collecting the permeate in a graduated cylinder for

a timed period.

6.3.3 Pressure Drop

Pressure drops for flow in the hollow Fiber tubings in the ultraflltration cartridge

was studied using a single hollow fiber tubing inserted inside a stopper-sealed glass column (Figure 6.2). The same length of the hollow fiber tubing as that in the membrane

module was used. The pressures at both inlet and outlet ends of the hollow fiber tubing were measured and used to determine the pressure drop in the tubing under various flow conditions. 176

r ~

Retentate

Permeate Ultraflltration hollow fiber module

Pump Xanthan gum solution

FIGURE 6.1: Schematic diagram of the ultraflltration system. pg) Glass Jacket (f^) Xanthan Retentate Broth Filtrate UF Membrane Q pump

FIGURE 6.2: Schematic diagrams of the experimental setup for pressure drop measurement. 178 6.3.4 Rheology Measurement

To study the rheological characteristics of xanthan solutions, shear stress and

apparent viscosity were measured at various shear rates. For high concentrations of

xanthan solutions, this was done with a Brookfield viscometer (RVTD II, spindle #5). For

low concentrations (up to 3.6%), a Rheometrics Asphalt Analyzer (RAA, Rheometrics

Inc., Piseataway, NJ) was used. All measurements were done at room temperature.

6.4 RESULTS AND DISCUSSION

6.4.1 Ultrafiltration

In membrane filtration, the volumetric flux of permeate, Jv, usually is modeled by

the following equation (Fane and Radovich, 1990):

j _APn-A x _ APm (6.1)

where APm is the transmembrane pressure drop, A k is the osmotic pressure difference

across the membrane, rj is the solution viscosity, and 27?/ is the total resistance of

filtration. In ultrafiltration of macromolecules such as xanthan polymer. An is negligible as compared to APm. The solution viscosity is also dependent on the concentration of xanthan gum and the shear rate (Lo et al„ 1995).

The ultrafiltration process needs to be optimized for several parameters, including xanthan concentration, solution viscosity, pumping (shear) rate, and trans-membrane pressure. As can be seen in equation (6.1), the filtrate flux, Jv, generally increases with increasing the transmembrane pressure, APm, and decreasing xanthan concentration and solution viscosity, rj. Since xanthan gum is a shear-thinning polymer, the viscosity of the 179 xanthan solution decreases with increasing shear rate or liquid velocity, and should result in an increase in the filtrate flux. Figure 6.3 shows the effects of xanthan concentration, APm,

and shear rate on filtrate flux.

The total resistance for filtration is highly dependent on pressure { APm), xanthan concentration (C), and shear rate (y) or flow rate ( Q ). Thus, equation (6.1) is not convenient to use in engineering design. However, the following empirical equations were found to be able to model the filtrate flux, J v (m3m-2s 1), as functions of C (g/L), APm

(Pa), and y (s'l);

Jv = 1.19 x 10'5 A P / 25 C 0 64 y °32 (6.2) where the proportional constant and exponent values were determined from the logarithmic plots of J v, vs. APm, C, and y, respectively (Figure 6.3). It is noted that the exponents associated with APm, y; and C all have an absolute value of less than 1, meaning that their effect on Jv is not linear. Equation. (6.2) is useful in predicting filtrate flux under various conditions, and thus, can be used in scale-up design.

6.4.2 Flow Behavior of Xanthan Solution

6.4.2.1 Rheoloeical characteristics

Xanthan gum solution is pseudoplastic and its rheological characteristics follow the power law equation:

T=K(-i^-r=~Kr (6.3) a r 180

100 A P = 25 psig

Shear Rate = 490 s

01 .1 1 10 100 Xanthan Concentration {%) (a) 10 2.5% Xanthan

Shear Rate = 490 s

1 1 i 1 o A P (105 Pa) (b) 10

1

A P = 27 psig

2.7% Xanthan 1 10 100 1000 Shear Rate (1/s) (c)

FIGURE 6.3: Determination of the empirical equation for UF process. 181

where t is shear stress, y is shear rate, K is consistency index, and n is flow behavior

index. The shear-thinning characteristics of xanthan solutions is shown in Figure 6.4. The apparent viscosity, r;, is defined as:

TJ = - = - K f ~ ' (6.4) y

Both K and n can be determined from the logarithmic plot of TJ vs. /show n in Figure 6.4.

For pseudoplastic fluid, n is less than 1. As the concentration of xanthan increases, K increases while n decreases (Figure 6.5). Both K and n approach a constant value at xanthan concentrations higher than 10%. At low xanthan concentration, n approaches 1 and K approaches the viscosity of water (1 cP). Similar results were reported by Galindo et a I. (1989), who studied the xanthan solution in the low concentration range (up to

2.5%).

The following empirical equations can be used to model n and K as the function of xanthan concentration in the range of 0.5% to 15.0% (wt/v):

n = 2.56 / C + 0.02 + 0.00095C - 0.0000033C2 (6.5)

K = -404 + 39.2C + 0.265C2 + 0.0006C3 (6.6)

6.4.2.2 Pressure drop for flow in tube

For ultrafiltralion of viscous xanthan solution, the liquid flow in the hollow fiber tubings was found to be laminar with a Reynolds number less than 10 (Chapter V).

Assuming no external forces (i.e., gravity), no entrance effect, incompressible fluid. 182

10000 Intercept = K Slope = n-

1000 3.6% 13.4% 10.3% 100 8.7% 1 .8 % 5.2%

10 0.9%

1

1 1 1 0 1000 Shear Rate (1/s)

FIGURE 6,4: Determination of K and n. K K (dyne s " /cm2 ) 10000 FIGURE 6.5: FIGURE 1000 100 10 1 1 0 ata Cnetain (g/L) Concentration Xanthan 090 30 Effect of xanthan concentration on on concentration xanthan of Effect 60 120 0.0 J 50 K 0.2 0.4 0.6 0.0 and and n. 183 184 laminar flow, and steady state, the flow for a power-law fluid in a cylindrical pipe as function of pressure drop, APt, is:

(6.7) where L is the length and r is the radius of the pipe. The average velocity, Uave, can be calculated from the volumetric flow rate, 0 , as follows:

Thus, the pressure drop for laminar flow in pipe can be estimated from Q as follows:

(6.9)

Both n and K in equation (6.9) are dependent on the concentration of xanthan gum in solution. The pressure drop thus will change with the xanthan concentration. As shown in Figure 6.6 for flow at Q = 16.8 mL/s and L = 40 cm, pressure drop inside the tube increased with increasing the xanthan concentration. Thus, more power input is required to maintain the same flow rate as the xanthan concentration increases. As also shown in this figure, the predictions from the model, equation (6.9), fit the date well.

Figure 6.7 shows the effect of pumping rates, Q, on the pressure drop for 52 g/L xanthan broth. As expected, the pressure drop increases with increasing the pumping rate.

However, the model predictions arc higher than actual measurements, especially at the lower pumping rates. FIGURE 6.6: FIGURE

Pressure Drop (psig) 100 Effect of xanthan concentration on pressure drop in UF module. UF in drop pressure on concentration xanthan of Effect 0 ata Cnetain (g/L) Concentration Xanthan 30 60 = 68 mL/s 16.86 = Q 0150 90 120 185 FiG VRE 6.7: VRE FiG Pressure Drop (psig) 20 25 30 10 15 o 5 0 Effect of pumping rate on pressure drop in UF module. UF in drop pressure on rate pumping of Effect OO 2 upn rt, (mL/s) Q rate, Pumping OO 4 .% ata Solution Xanthan 5.2% 10 6 8 12 186 187

6.4.3 Energy Consumption for Ultrafiltration

6.4.3.1 Power consumption for pumping

The major energy consumption in the ultrafiltration process is the power consumed for pumping the fluid in the hollow fiber tubings. Neglecting potential energy and kinetic energy changes for flow in horizontal pipes, the required shaft work, Ws> from the pump is equal to AS*/p. The power consumption, Pp, for the process is:

PP = {pQ)Ws =QAP, (6.10)

6.4.3.2 Energy consumption

The energy requirement in concentrating the xanthan fermentation broth by ultrafiltration is mainly attributed to the pumping. The energy consumption per unit volume of water removed from ultrafiltration can be estimated from the power consumption divided by the filtrate rate, F\ i.e. P/F.

Figure 6.8 shows filtrate rate, power consumption, and energy consumption as affected by the xanthan concentration. As expected, more energy is required to remove water from a higher concentration xanthan solution.

6.4.3.3 Effect of shear (pumping) rate

In general, a sufficiently high pumping (shear) rate should be used in xanthan ultrafiltration; otherwise, the solution viscosity may become formidably high for the process. However, the energy requirement per unit volume of water removed from ultrafiltration could be higher at higher pumping rates, even though the filtrate rate is also higher. This is because power consumption increases more proportionally but filtrate rate increases less proportionally with increasing the pumping rate. Therefore, there is an FIGURE 6.8: FIGURE

Effect of xanthan concentration on filtrate rate, power consumption, power rate, filtrate on concentration xanthan of Effect Filtrate Rate, F (mL/s) 0.2 0.4 - - 0.4 0.3 - - 0.3 0.1 5r r .5 0 - 0 and energy consumption. energy and - - *« u 0k & 'w' £ E o v O G O G VI s CL, it b. 0 2 0 4 3 4 50 5 1 0 = 66 mL/s 16.66 = Q A - 2 psig 27 - P ata Cnetain (g/L) Concentration Xanthan 20 060 40 80 P/F 100 20 0 3

P/F (J/mL) 188 189 optimal pumping rate for minimum energy consumption during ultrafiltration, as is shown in Figure 6.9. The optimal pumping rate for the process, however, should be determined by also taking into account the membrane replacement cost, which decreases with increasing the pumping rate because of the improved filtrate flux and reduced membrane area required for the process.

6.4.3.4 Effect of transmembrane pressure

In general, the energy requirement for ultrafiltration should decrease with increasing APm because of increased filtrate rate at higher APm. Therefore, ultrafiltration should be performed at the highest possible pressure. For the hollow fiber cartridge used in this study, the maximum operable pressure was only -30 psig. However, for industrial operation using spiral wound membrane units, the pressure can go up to 70 psig. Figure

6.10 shows that energy consumption decreases as increasing the transmembrane pressure.

Power consumption was assumed not being significantly affected by the transmembrane pressure.

6.4.4 Process Economics

6.4.4.1 Operating costs for ultrafiltration

Only energy (pumping) and membrane costs were considered in estimating the operating costs for ultrafiltration to concentrate xanthan solution from 2.5% to 15%. The energy cost for pumping was estimated at the rate of $0.05/KWh at a pumping efficiency of

70%. The cost of membrane was estimated at $200/m2 as quoted by the manufacturer,

KOCH Membrane Co. (Gach, 1995). The operating life of the membrane should be about one year. Thus, the membrane replacement cost is $200/m2 per year. FIGURE 6.9: FIGURE

Power Consumption, P (W) S 0.4 OS a a A SB Effect of pumping rate on filtrate rate, power consumption, power rate, filtrate on rate pumping of Effect 0.0 0.2 0.6 0.8 0 and energy consumption. energy and . 2.7% Xanthan at at Xanthan 2.7% . 2 upn Rt, (mL/s) Q Rate, Pumping 4 6 1 1 1 1 18 16 14 12 10 8 A P' = 27 psig 27 = P' P/F 0

P/F (J/mL) 190 FIGURE 6.10: FIGURE

Power Consumption, P (W) 0.3 * 0.4 £ a 0-2 2 <3 S Effect of transmembrane pressure on filtrate rate, power rate, filtrate on pressure transmembrane of Effect 0.0 0.5 0.7 0.8 0.6 0 consumption, and energy consumption. energy and consumption, rnmmrn Pesr (psig) Pressure Transmembrane .% Xnhn t = 0 mL/s 10 = Q at Xanthan 2.7% 20 40 P/F 60 80 0

P/F (J/mL) 191 192

The operating costs are affected by the operating conditions, e.g., shear rate and

transmembrane pressure. As shown in Figure 6.11, the optimal shear rate for minimum

operating costs may shift greatly if the membrane cost changes. In general, a higher shear

should be used to offset increase in the membrane cost. Figure 6.12 shows the effect of APm on the operating costs. It is noted that the possible effects of shear and pressure on

membrane fouling and life were not considered in these estimation. At the optimal conditions ( APm ~ 70 psig, shear rate = 153 s 1) with the quoted membrane cost of

$200/m2, it costs -$1.5 per kg xanthan gum processed by ultrafiltration to remove -33.3

liters water from the 2.5% xanthan fermentation broth.

6.4.4.2 Energy costs for alcohol distillation

The energy used to recover the ethanol used in precipitating xanthan gum from the

fermentation broth was estimated. Two volumes of alcohol (99% v/v) are used to each

volume of xanthan broth (2.5% w/v). Thus, it will need to separate the 65% (v/v) ethanol

from water to 99% ethanol by distillation. Based on a typical distillation unit with 40 stages, 1.5 reflux ratio, and 453.6 kg/h (1000 lbm/h) feed rate, the energy input for

distillation as calculated by using ASPEN is 925,192 KJ/kg xanthan produced. Assuming that the steam used in distillation has a heating capacity of 1024 KJ/lbm and costs $3.0 per

1000 lbm, the energy cost for alcohol precipitation, without considering any alcohol loss in

the process, is estimated at $2.25 per kg xanthan gum processed.

6.4.4.3 Total xanthan recovery costs

It is clear that ultrafiltration can be used economically to concentrate xanthan broth

before alcohol precipitation. Large amounts of energy can be saved by using ultrafiltration

as it requires no vaporization in removing water. The total energy consumption for the recovery process using ultrafiltration followed by alcohol precipitation is shown in Figure 193

3 c A Pm = 70 psig A 6 e Membrane: $ 200 / m a X 2 Total M J4

1 Membrane

u Energy

o o 100 200 300 400 500 (a) 4 A Pm = 70 psig s Total A Membrane: $ 400 / m 3 A X M .M 2 Membrane

S Energy

0 o 1 oo 200 300 400 500 Shear Rate (1/s) (b)

FIGURE 6.11: Operating costs for UF at transmembrane pressure equals 70 psig under various shear rate with: (a) membrane cost at $200/m2; and (b) membrane cost at $400/m2. 194

3 e Shear Rate = 153 s « Total JZ4-t Membrane: $ 200 / m ea(3 X 2 M l

Membrane Vi 1 CL oV) U Energy O- 0 0 10 20 3040 50

A P (104 Pa)

FIGURE 6.12: Operating costs for UF at shear rate equals 153 s 1 under various transmembrane pressure with membrane cost at $200/m2. 195

6.13. It is clear that large amounts of energy can be saved if ultrafiltration is used to

remove water before alcohol precipitation. For the range of 2.5% to 15% xanthan, more

energy can be saved by concentrating to a higher xanthan concentration with ultrafiltration.

However, the optimal xanthan concentration from ultrafiltration should be based on the

total operating costs, which also include membrane cost

6.5 CONCLUSIONS

Ultrafiltration can be economically used to concentrate xanthan broth and to reduce

the amounts of alcohol and energy used in the xanthan recovery process by -80%. The

process performance and economics are affected by the transmembrane pressure, shear

rate, and xanthan concentration. The models developed for ultrafiltration of xanthan

solution and pseudoplastic fluid flow in tube fit the experimental data well and can be used

in process analysis.

6.6 NOTATION

C Xanthan concentration in solution (g/L)

-F'iw Energy loss due to fluid friction between wall and fluid (J)

J v Volumetric permeate flux for solution ultrafiltration (m3 nr2 s 1)

K The consistency index (dyn-sn/cm2)

L The length of the tube (m) n The flow behavior index (-) Power (J s’1, W) P p Transmembrane pressure defined by equation (6.1) (Pa, psig) I E 6.13: RE U FIG

Energy (kj / kg Xanthan) Total energy consumption for xanthan gum recovery process recovery gum xanthan for consumption energy Total 1000000 200000 400000 600000 800000 using UF followed by alcohol precipitation. alcohol by followed UF using 25 Distillation ia Xnhn ocnrto (g/L) Concentration Xanthan Final Total 50 75 100 UF x 500 x UF 2 150 125 1% 197 AP( Pressure drop ai the end of the tube (Pa, psig)

Q Pumping rate (mL/s) /?, Resistance owing to a specific source (Pa nr3 m2-s) r Radius of a tube (m) Uz Instantaneous velocity vector in z direction (m s1)

UZi ave Mean velocity in z direction (m s 1)

UL^^^Maximum velocity in z direction (m s1)

Ws Shaft work done by the pump of the system (J)

Greek letters y Shear rate at wall (s-1) tj Viscosity (cP) k Ratio of circumference of a circle to its diameter (3,141 592 65...)

Ap Osmotic pressure difference (Pa) t Shear stress (N n r2, kg m 1 s*2)

6.7 REFERENCES

Barker, P. E. & Till, A. (1992). Using multistage techniques to improve diafiltration fractionation efficiency. J. Membrane Sci., 72, 1-11.

Brodkey, R. S. & Hershey, H. C. (1988). Transport Phenomena: A Unified Approach. McGraw-Hill, New York.

Cottrell, I. W., Kang, K. S. & Kovacs, P. (1980). Xanthan Gum. In Handbook of Water- Soluble Gums and Resins. McGraw-Hill, New York, Chapter 24.

Dintzis, F. R., Babcock, G. E. & Tobin, R. (1970). Studies on dilute solutions and dispersions of the polysaccharide from Xanthomonas campestris NRRL B-1459. Carbohyd. Res., 13, 257-267.

Flahive III, J. J., Foufopoulos, A. & Etzel, M. R. (1994). Alcohol Precipitation of Xanthan Gum from Pure Solutions and Fermentation Broths. Separation Sci. Technol ., 29(13), 1673-1687. 198 Funahashi, H., Yoshida, T. & Taguchi, H. (1989). Effect of Shear Stress on Xanthan Gum Production. In Bioproducts and Bioprocesses. Springer-Verlag, Berlin, pp. 359-370.

Gach, G. (1995) KOCH Membrane Co., Ann Arbor, MI., personal consultation.

Galindo, E., Torrestiana, B. & Garcia-Rejdn, A. (1989). Rheological characterization of xanthan fermentation broths and their reconstituted solutions. Bioprocess Eng. , 4, 113- 118.

Garcfa-Ochoa, F. & Casas, J. A. (1994). Apparent yield stress in xanthan gum solutions at low concentrations. The Ckem. Eng. / , 53, B41-B46.

Garcfa-Ochoa, F., Casas, J. A. & Mohedano, A. F. (1993). Xanthan Precipitation from Solutions and Fermentation Broths. Sep. Sci. Technol, 28(6), 1303-1313.

Gonzales, R„ Johns, M. R„ Greenfield, P. F. & Pace, G. W. (1989). Xanthan gum precipitation using ethanol. Process Biochemistry , 24(6), 200-203.

Hannote, M., Flores, F„ Torres, L. & Galindo, E. (1991). Apparent yield stress estimation in xanthan gum solutions and fermentation broths using a low-cost viscometer. The Chem. Eng. J., 45, B49-B56.

Holzwarth, G. (1978). Molecular weight of xanthan polysaccharide. Carbohyd. Res., 66, 173-186.

Jonsson, A.-S. (1993). Influence of shear rate on the flux during ultrafiltration of colloidal substances. J. Membrane Sci., 79, 93-99.

Kang, K. S. & Pettitt, D. J. (1993). Xanthan, Gellan, Welan, and Rhamsan. In Industrial Gums; Polysaccharides and Their Derivatives. Academic Press, San Diego, CA, pp. 341 - 371.

Lim, T., Uhl, J. T. & Prud’homme, R. K. (1984). Rheology of self-associating concentrated xanthan solutions. J. Rheology, 28(4), 367-379.

Lo, Y. M., Yang, S. T. & Min, D. B. (1995). Xanthan gum recovery from fermentation broth using ultrafiltration. J. Membrane Sci. (submitted).

Margaritis, A. & Pace, G. W. (1985). Microbial Polysaccharides. In Comprehensive Biotechnology, vol. 3. Pergamon Press, New York, pp. 1005-1043

McNeil, B. & Harvey, L. M. (1993). Viscous Fermentation Products. Critical Reviews in Biotechnol., 13(4), 275-304.

Meyer, E. L., Fuller, G. G., Clark, R. C. & Kulicke, W. M. (1993). Investigation of xanthan gum solution behavior under shear flow using rheooptical techniques. Macromolecules, 26, 504-511.

Milas, M., Rinaudo, M. & Tinland, B. (1985). The viscosity dependence on concentration, molecular weight and shear rate of xanthan solutions. Polymer Bulletin, 14, 157-164. 199 Mulder, M. (1991). Basic principles of membrane technology, Ch. 6. Kluwcr Academic Publishers, New York, pp. 198-212.

Norton, I. T., Goodall, D. M., Frangon, S. A., Morris, E. R. & Rees, D. A. (1984). Mechanism and dynamics of conformational ordering in xanthan polysaccharide. J. Mol. Biol. 175, 371-394.

Oviatt, Jr., H. W. & Brant, D. A. (1994). Voscoelastic behavior of thermally treated aqueous xnathan solutions in the semidilute concentration regime. Macromolecules, 27, 2402-2408.

Pickering, K. D. & Wiesner, M. R. (1993). Cost model for low-pressure membrane filtration. J. Environ. Eng., 119(5), 772-797.

Rochefort, W. E. & Middleman, S. (1987). Rheology of xanthan gum; salt, temperature, and strain effects in oscillatory and steady shear experiments. J. Rheology, 31(4), 337- 369.

Rogovin, S. P., Albrecht, W. J. & Sohns, V. (1965). Production of industrial grade polysaccharide B-1459. Biotechnol. B'toeng., 7. 161-169.

Shiloach, J. & Kaufman, J. B. (1986). Hollow fiber microfiltration methods for recovery of rate Basophilic Leukemia cells from tissue culture media. Biotechnol . Prog., 2(4), 230- 233.

Swift, C. J. & Atkins, W. S. (1983). Alcohol recovery by distillation. M onograph: European Brewery Convention, 185-202.

Tam, K. C. & Tiu, C. (1993). Rheology of water-soluble polymers: a comparative study on the effect of monovalent salt. Polynu-Plast. Technol. Eng., 32(1 &2), 123-138.

Torres, L. G., Brito, E., Galindo, E. & Choplin, L. (1993). Viscous behaviour of xanthan aqueous solutions from a variant strain of Xanthomonas campestris. J. Fermen . Bioeng., 75(1), 58-64.

Tweddle, T. A., Striez, C., Tam, C. M. & Hazlett, J. D. (1992). Polysulfone membranes I. Performance comparison of commercially available ultrafiltration membranes. Desalination, 86, 27-41. van Oers, C. W., Vorstman, M. A. G., Muijselaar, W. G. H. M. & Kerkhof, P. J. A. M. (1992). Unsteady-state flux behavior in relation to the presence of a gel layer. J. Membrane Sci., 73, 231-246.

Yeh, H.-M. & Cheng, T.-W. (1993). Osmotic-pressure model with permeability analysis for ultrafiltration in hollow-fiber membrane modules. Separation Technol., 3, 91-98. CHAPTER VII

CONCLUSIONS AND RECOMMENDA TIONS

7.1. Centrifugal, Packed-Bed Bioreactor for Xanthan Gum Production

The centrifugal, packed-bed bioreactor was able to produce xanthan gum at a

productivity twice of that for the present industrial process. The intimate air, liquid, and

cell contacts achieved via passing liquid medium and air through the porous fibrous matrix

achieved an enhanced oxygen transfer rate and high xanthan productivity. Separation of

cell growth and xanthan production into two stages also improved xanthan fermentation

with enhanced xanthan production rate and yield. Hie cells immobilized in the bioreactor

was able to be repeatedly used for xanthan fermentation to achieve continuous production

of cell-free xanthan broth. Therefore, subsequent cell-removal from fermentation broth is

eliminated. The production of cell-free fermentation broth would benefit the downstream

ultrafiltration process by eliminating membrane fouling caused by cells and cell debris

(DNA, RNA, etc.) that would otherwise be present in the xanthan broth.

The oxygen transfer rate in CPBR-GC was the highest among various systems

studied at 3.5% xanthan concentration. Therefore, the CPBR system is the preferred system for xanthan gum fermentation in terms of oxygen transfer ability. At high xanthan concentrations, the ability of aeration by conventional agitation decreased. The direct air-

200 201 liquid contact provided in the rotational fibrous matrix thus becomes an important mechanism for aeration in CPBR.

It is possible to further increase the xanthan productivity by increasing the rotational speed of the fibrous matrix and recirculation rate. The immobilized cell density in the reactor can be increased to further increase the reactor productivity. However, this would require a parallel increase in oxygen transfer rate. Because it is more important to use a high rotational speed for high concentration xanthan solutions, a better constructed rotating fibrous matrix holder that can be operated at a higher speed than that has been achieved in this study will be needed to provide better oxygen transfer. Increase in the flow rate in the recirculation stream, which provided direct contacts of gas, liquid, and matrix (cells) in

CPBR, should be able to further enhance the oxygen transfer rate and xanthan productivity in the reactor. Since oxygen transfer in CPBR is not as sensitive to increasing in the xanthan concentration as the conventional stirred tank reactor, it is possible to produce xanthan at a concentration higher than 3.5% achieved in this study.

Furthermore, it is noted that the rotation matrix in the centrifugal bioreactor need not be an internal part of a fermentor. It can be used as an external mass transfer device for existing conventional fermentors. This CPBR can also be operated with continuous feed for xanthan production.

7.2 Ultrafiltration of Xanthan Gum Fermentation Broth

Ultrafiltration with the hollow fiber cartridge can concentrate dilute xanthan fermentation broth (-2.5%) to a concentration of -13% or higher, with a recovery yield of greater than 95%. The process is stable and does not harm the qualities of the xanthan 202 polymer. The concentrated xanthan broth would require proportionally smaller amounts of alcohol for xanthan gum production. With a fivefold increase in (he xanthan concentration, the amounts of alcohol needed in xanthan recovery would be reduced by 80%. Also, water and solutes present in the permeate from ultrafiltration may be recycled for use in fermentation. Compared to the present alcohol precipitation method for xanthan gum recovery, the ultrafiltration process could reduce the energy costs by over 80%, and is more economical and environmentally friendly.

Further research should be focused on methods that can greatly enhance filtrate flux. Electro-ultrafiltration and high-shear ultrafiltration processes should have great potentials to achieve this goal. It is also recommended to test the process with spiral- wound ultrafiltration module at an elevated transmembrane pressure (i.e. 70 psig) and to test the long-term performance of the ultrafiltration membrane. BIBLIOGRAPHY

Ahmad, M.N.; Holland, C.R.; McKay, G. Mass transfer studies in batch fermentation: mixing characteristics. J. Food Eng. 1994, 23, 145-158.

Anderson, J.G.; Blain, J.A.; Divers, M.; Todd, J.R. Use of the disc fermentor to examine production of citric acid by Aspergillus niger. Biotechnol. Letters , 1980, 2, 99.

Atkinson, B. Immobilized cells, their applications and potential. In Process Engineering Aspects of Immobilized Cell Systems. Webb, C.; Black, G.M.; Atkinson, B. (Eds.), Pergamon Press, New York. Ch. 1, 1986.

Atkinson, B.; Cunningham, J.D.; Pinches, A. The biomass hold-ups and overall rates of substrate (glucose) uptake of support particles containing a mixed microbial culture. Chem. Eng. Res. Des. 1984, 62, 155.

Bailey, J.; Ollis, D.E. Biochemical Engineering Fundamentals. 2nd Ed., McGraw Hill, New York, NY. 1986.

Barker, P. E.; Till, A. Using multistage techniques to improve diafiltration fractionation efficiency. J. Membrane Sci. 1992,72, 1-11.

Behrens, U.; Klima, M.; Fiedler, S. Growth and accumulation of polysaccharide by Xanthomonas campestris. Z Allg. Mikrobiol. 1980,20, 209-213.

Bezemer, L.; Ubbink, J.B.; de Kooker, J.A.; Kuil, M.E.; Leyte, J.C. On the conformational transitions of native xanthan. Macromolecules, 1993,26, 6436-6446.

Bhavaraju, S.M.; Russel, T.W.F.; Blanch, H.W. The design of gas sparged devices for viscous liquid systems. AIChE J. 1978,24(3), 454-466.

Bier, M. Electrokinetic membrane processes. In Membrane Processes in Industry and Biomedicine. 1971. pp 233-266.

Bilanovic, D.; Shelef, G.; Green, M. Xanthan fermentation of citrus waste. Bioresource Technol., 1994, 48, 169-172.

Black, G.M. Characteristics and performance of immobilized cell reactors. In Process Engineering Aspects of Immobilized Cell Systems. Webb, C.; Black, G.M.; Atkinson, B. (Eds.), Pergamon Press, New York. Ch. 5, 1986.

Blain, J.A.; Anderson, J.G.; Todd, J.R.; Divers, M. Cultivation of filamentous fungi in the disc fermentor. Biotechnol . Letters, 1979,1, 267.

203 204

Bordi, F.; Cametti, C.; Paradossi, G. Radiowave dielectric properties of xanthan in aqueous solutions. J. Phys. Chem., 1995, 99, 274-284.

Bowen, W.R.; Sabuni, H.A.M. Pulsed electrokinetic cleaning of cellulose nitrate microfiltration membranes. Ind Eng. Chem. Res. 1992, 31(2), 525-533.

Brodkey, R. S. The Phenomena of Fluid Motions. Hoelscher, H. (Ed.), 1967, Robert S. Brodkey, Columbus, OH. Ch. 15.

Brodkey, R.S.; Hershey, H.C. Transport Phenomena: A Unified Approach. Clark, B.J. (Ed.), 1988, McGraw-Hill, New York, NY. Ch. 9.

Brooker, B.E.; Fuller, R. In Microbial Ultrastructure: The Use of The Electron Microscope. Fuller, R.; Lovelock, D.W. (Eds.), 1976. Academic Press, New York, NY. p. 87.

Bruin, S.; Kikkert, A.; Weldring, J.A.G.; Hiddink, J. Overview of concentration polarization in ultrafiltration. Desalination, 1980, 35, 223-242.

Buchholz, K. (Ed.) Characerization of immobilized biocatalysis. In DECHEMA monographs, Vol. 84, 1979. Verlag Chemie, New York, NY. pp 1724-1731.

Bukharov, A.V.; Skvortsov, I.M.; Ignatov, V.V.; Shashkov, A.S.; Knirel, Y.A.; Dabrowski, J. Structure of the O-specific Polysaccharide of Xanthomonas campestris NCPPB 45 Lipopolysaccharide. Carbohydrate Res., 1993, 24J, 309-316.

Burns, R.G. Microorganisms and soil surfaces. In Adhesion of microorganisms to surfaces. Ellwood, D.C.; Melling, J.; Rutter, P. (Eds.), 1979. Academic Press, New York, NY. p. 109.

Cadmus, M.C.; Rogovin, S.P.; Burton, K.A.; Pittsly, J.E.; Knutson, C.A.; Jeanes, A. Colonial variation in Xanthomonas campestris NRRL B-1459 and characterization of the polysaccharide from a variant strain. Can. J. Microbiol. 1976,22, 942-948.

Cantor, C.R.; Schimmel, P.R. Biophysical chemistry, part II: techniques for the study of biological structure and function. W.H. Freeman and Company. New York, NY. 1980, pp 643-685, Ch. 12.

Charles, M. Technical aspects of the rheological properties of microbial cultures. Adv. Biochem. Eng. 1978, 8, 1-62.

Cheryan, M. Ultrafiltration handbook. Technomic, Lancaster, PA, 1986.

Chowdia, P.; Wasan, D.T.; Gidaspow, D. Electrokinetic phenomena in the filtration of colloidal particals suspended in nonaqueous media. AIChEJ. 1981, 27(6), 975-983.

Christensen, B.E.; Knudsen, K.D.; Smidsrod, O.; Kitamura, S.; Takeo, K. Temperature- induced conformational transition in xanthans with partially hydrolyzed side chains. Biopolymers, 1993, 33, 151-161. 205 Christensen, B.E.; Smidsr 0d, O.; Elgsaeter, A.; Strokke, B.T. Depolymerization of double-stranded xanthan by acid hydrolysis: characterization of partially degraded double strands and single-stranded oligomers released from the ordered structures. Macromolecules, 1993, 26, 6111-6120.

Comberbach, D.M.; Bu'lock, J. Continuous ethanol production in the gas-lift tower fermentor. Biotechnol. Letters, 1984,6, 129.

Costerton, J.W.; Marks, I. Localization of enzymes in prokaryotic cells. In Electron Microscopy o f Enzymes: Principles and Methods. 1977, vol. 5, p. 99.

Cotrell, I.W.; Kang, K.S.; Kovacs, P. Xanthan gum. In Handbook of Water-Soluble Gums and Resins', Davidson, R. L. (Ed.), 1980; McGraw-Hill: New York; Chapter 24.

Cottrell, G.A.; Kang, K.S. Xanthan gum, a unique bacterial polysaccharide for food applications. Dev. Ind. Microbiol. 1978, 19, 117-131.

Crueger, A.; Crueger, W. Carbohydrates. Biotechnology, 1984, 6a, 421-457.

Cysewski, G.R.; Wilke, C.R. Fermentation kinetics and process economics for the production of ethanol. Biotechnol. Bioeng. 1977, 19, 1125. daFonseca, M.M., Black, G.M.; Webb, C. Reactor configurations for immobilized cells, in Process Engineering Aspects of Immobilized Cell Systems. Webb, C.; Black, G.M.; Atkinson, B. (Eds.), Pergamon Press, New York. Ch. 4, 1986.

Davidson, I. W. Production of polysaccharide by Xanthomonas campestris in continuous culture. FEMS Microb. Lett. 1978, 3, 347-349.

Deavin, L.; Jarman T.R.; Lawson, C.J.; Righelato, R.C.; Slocombe, S. The production of alginic acid by Azotobacter vinelandii in batch and continuous culture. ACS Symp. Ser. 1977, 45, 14-26.

DeVuyst, L,; VanLoo, J.; Vandamme, E.J. Two-step fermentation process for improved xanthan production by Xanthomonas campestris NRRL B-1459. J. Chem. Tech. Biotechnol. 1987, 39, 263-273.

DiCosmo, F.; Tanaka, H.; Neumann, A.W. Cell immobilization by adsorption to glass fibre mats. In Immobilized Biosystems: Theory and Practical Applications. Veliky, I.A.; McLean, R.J.C. (Eds.), 1994. Blackie Academic & Professional, Glasgrow, UK.

Dintzis, F. R.; Babcock, G. E.; Tobin, R. Studies on dilute solutions and dispersions of the polysaccharide from Xanthomonas campestris NRRL B-1459. Carbohyd. Res . 1970, 13, 257-267.

Drozd, J.W.; Rye, A.J. Method for improving the filterability of a microbial broth. U.S. Patent 4,729,958; 1988.

Eisenberg, D.; Crothers, D. Physical chemistry: with applications to the life sciences. The Benjamin/Cummings Publishing, Menlo Park, CA. 1979, pp 700-741, Ch. 15. 206 Engelskirchen, K.; Stein, W.; Bahn, M.; Schieferstein, L.; Schindler, L.; Schmid, R. Process for producing Xanthomonas biopolymers. U.S. Patent 4,618,582. October 21, 1986.

Fane, A.G. Ultrafiltration: Factors influencing flux and rejection. In Progress in Filtration and Separation, Vol. 4. Wakeman, R.J. (Ed.), 1986. Elsevier Scientific Publishing, Amsterdam.

Fane, A.G.; Radovich, J.M. Membrane systems. In Separation Processes in Biotechnology. Asenjo, J.A. (Ed.), 1990; M. Dekker, New York, NY; Ch. 8 .

Fiedler, S. and Behrens, U. Growth of Xanthomonas campestris and xanthan accumulation. ZeitschriftfUr allgemeine Mikrobiologie 1984, 24 (4): 223-230.

Flahive III, J.J.; Foufopoulos, A.; Etzel, M.R. Alcohol precipitation of xanthan gum from pure solutions and fermentation broths. Separation Sci. Technol. 1994, 29(13), 1673- 1687.

Fletcher, M.; Floodgate, G.D. In Microbial Ultrastructure: The Use of The Electron Microscope. Fuller, R.; Lovelock, D.W. (Eds.), 1976. Academic Press, New York, NY. p. 101.

Flores, F.; Torres, L.G.; Galindo, E. Effect of the dissolved oxygen tension during cultivation of X. campestris on the production and quality of xanthan gum. J. Botechnol. 1994, 34, 165-173.

Foerberg, C.; Enfors, S.O.; Haeggstroem, L. Control of immobilized, nongrowing cells for continuous production of metabolites. Eur. J. Appl. Microbiol. Biotechnol. 1983, 17, 143.

Fowler R.; Balasundaram, V. De-aeration under the action of centrifugal force. Proceedings of Water Management Offshore, Aberdeen, Scotland, 1989.

Fricker, R. European Brewery Convention, Monograph - IX, 170, 1983.

Funahashi, H.; Maehara, M.; Taguchi, H.; Yoshida, T. Effects of agitation by flat-bladed turbine impeller on microbial production of xanthan gum. Journal of Chemical Engineering o f Japan, 1987,20(7), 16-22.

Galindo, E.; Nicnow, A.W. Mixing of highly viscous simulated xanthan fermentation broths with the Lightnin A-315 impeller. Biotechnol. Prog. 1992, 8, 233-239.

Galindo, E.; Nicnow, A.W. Performance of the Scaba 6 SRGT agitator in mixing of simulated xanthan gum broths. Chem. Eng. Technol., 1993, 16, 102-108.

Galindo, E.; Torrestiana, B.; Cuernavaca; Garcia-Rej 6n, A. Rheological characterization of xanthan fermentation broths and their reconstituted solutions. Bioprocess Eng. 1989, 4, 113-118.

Gamini, A.; Mandel, M. Physicochemical properties of aqueous xanthan solutions: static light scattering. Biopolymers , 1994, 34, 783-797. 207 Garcfa-Ochoa, F.; Casas, J.A. Apparent yield stress in xanthan gum solutions at low concentrations. The Chem. Eng. J., 1994, 53, B41-B46.

Garcfa-Ochoa, F.; Casas, J.A.; Mohedano, A.F, Xanthan precipitation from solutions and fermentation broths. Sep. Sci. Technol., 1993 ,28(6), 1303-1313.

Gill, W.N.; Wiley, D.E.; Fell, C.J.D.; Fane, A.G. Effect of viscosity on concentration polarization in ultrafiltration. AfChEJ. 1988 ,34(9), 1563-1567.

Gonzales, R.; Johns, M. R.; Greenfield, P. F,; Pace, G. W. Xanthan gum precipitation using ethanol. Process Biochemistry, 1989, 24(6), 200-203.

Green, A.A.; Hughes, W.L. In Methods in Enzymology. Colowick, S.P.; Kaplan, N.O. (Eds.), 1955. Academic Press, New York, NY. Vol. 1, p. 67.

Hannote, M.; Flores, F.; Torres, L.; Galindo, E. Apparent yield stress estimation in xanthan gum solutions and fermentation broths using a low-cost viscometer. The Chemical Engineering Journal, 1991, 45, B49-B56.

Haze, A.; Kimura, K.; Hashimoto, K. Method of dehydrating macromolecular substances. European Patent 0343459A2.

Henry Jr., J.D.; Lawler, L.F.; Kuo, K.H.A. A solid/liquid separation process based on cross flow and electrofiltration. AIChE J. 1977, 23(6), 851-859.

Henzler, H.J. Engineering design data for stirred vessels as fermentors. Chem. Ing. Tech. 1982, 54(5), 461-476.

Herbst, H.; Peters, H.-U.; Suh, I.-S.; Schumpe, A.; Deckwer, W.-D. Oxygen transfer during xanthan fermentation. 1987 AIChE Annual Meeting, 175a; New York, 1987.

Herbst, H.; Schumpe, A.; Deckwer, W.-D. Xanthan production in stirred tank fermentors: oxygen transfer and scale-up. Chem. Eng. Technol. 1992, 15, 425-434.

Herbst, H.; Suh, I. S.; Peters, H. U.; Schumpe, A.; Deckwer, W. D. Comparison of various fermentor types used for production of xanthan. DECHEMA Biotechnol. Conference. 1989, 3, 495-498.

Himmelsbach, W. Studies on mass and heat transfer and on homogenization in stirred fermentors using xanthan as an example, Chem. Ing. Tech., 1985 ,57(6), 548-549.

Ho, C.S.; Ju, L.K.; Baddour, R.F. The anomaly of oxygen diffusion in aqueous xanthan solutions. Biotechnol. Bioeng., 1988, 32, 8-17.

Holzwarth, G. Molecular weight of xanthan polysaccharide. Carbohyd. Res. 1978, 66, 173-186.

Hong, J.; Lee, C.K. Membrane separation coupled with electrophoresis. In: Biochemical Engineering IV. New York, 131-144. 208 Ielpi, L.; Couso, R.O.; Dankert, M.A. Xanthan gum biosynthesis: pyruvic acid acetal residues are transferred from phosphoenolpyruvate to penta-saccharide-P-P-lipid. Biochem. Biophy. res. comm. 1981, 120, 1400-1407.

Inatomi, S.; Jinbo, Y.; Sato, T.; Teramoto, A. Isotropic-liquid crystal phase equilibrium in semiflexible polymer solutions: effects of molecular weight and ionic strength in polyelectrolyte solutions. Macromolecules, 1992, 25, 5013-5019.

Inloes, D.S.; Smith, W.J.; Taylor, D.P.; Cohen, S.N.; Michaels, A.S.; Robertson, C.R. Hollow-fiber membrane bioreactor using immobilized E. coli for protein synthesis. Biotechnol. Bioeng., 1983, 25, 2653.

Jannsson, P. K.; Kcnne, L.; Lindberg, B. Structure of the extracellular polysaccharide from Xanthomonas campestris. Carbon. Research , 1975, 45, 275-282.

Jarman, T.R. Xanthan gum. Ear. Pat. Appl. EP66,377. 1983.

Jarman, T.R.; Deavin, L.; Slocombe, S.; Righelato, R.C. Investigation of the effect of environmental conditions on the rate of exopolysaccharide synthesis in Azotobacter vinelandii. J. Gen . Microbiol. 1978, 107(1), 59-64.

Jarman, T.R.; Pace, G.W. Energy requirements for microbial exopolysaccharide synthesis. Arch. Microbiol. 1984, 137, 231-235.

Jeanes, A. Applications of extracellular microbial polysaccharide-polyelectrolytes: Review of literature, including patents. J. Polym. Sci. Symp. 1974, 45, 209-227.

Jeanes, A.; Rogovin, S.P.; Cadmus, M.C.; Silman, R.W.; Knutson, C.A. Agric. Res. Service, U.S. Dept, of Agric. ARS-NC-51, 1976.

Jeanes, A.R.; Pittsley, J.E.; Senti, F.R. Polysaccharide B-1459: hydrocolloid polyelectrolytes produced from glucose by bacterial fermentation. J. Appl. Polymer Sci. 1961,5,519-526.

Jdnsson, A.-S. Influence of shear rate on the flux during ultrafiltration of colloidal substances. J. Membrane Sci. 1993, 79, 93-99.

Ju, L.-K.; Zhao, S. Xanthan fermentations in water/oil dispersions. Biotechnol. Techniques 1993, 7(7), 463-468.

Kan, J.K.; Shuler, M.L. Urocanic acid production using whole cells immobilized in a hollow fiber reactor. Biotechnol. Bioeng., 20: 217, 1978.

Kang, K. S.; Cottrell, I. W. Polysaccharides. In Microbial Technology, vol. I, Microbial Process', Peppier, H. J.; Perlman, D. (Eds.), 1979; Academic Press, New York; Chapter 13.

Kang, K.S.; Burnett, D.B. Increasing the viscosity of aqueous solutions of xanthan gum. Fr. Demande 2,318,926. 1977. 209 Kang, K.S.; Pettitt, D.J. Xanthan, Gellan, Welan, and Rhamsan. In Industrial Gums: Polysaccharides and Their Derivatives ; Whistler, R. L,; BeMiller, J. N. (Eds.), 1993; Academic Press, San Diego, CA; pp341-371.

Karamanev, D.G.; Nagamune, T.; Endo, I. Hydrodynamic and Mass Transfer Study of a Gas-Liquid-Solid Draft Tube Spouted Bed Bioreactor. Chem. Eng . Sci. 1992,47(13), 3581-3588.

Kawase, Y.; Tsujimura, M. Enhancement of oxygen transfer in highly viscous non- Newtonian fermentation broths. Biotechnol. Bioeng. 1994,44, 1115-1121.

Kennedy, J. F.; Bradshaw, I. J. Production, properties and applications of xanthan. In Progress in Industrial Microbiology, vol. 19. Bushell, M.E. (Ed.), 1984; Elsevier Science, New York, NY; pp. 319-371.

Kennedy, J.F.; Jones, P.; Barker, S.A.; Banks, G.T. Factors affecting microbial growth and polysaccharide pioduction during the fermentation of Xanthomonas campestris cultures. Enzyme Microb. Technol. 1982, 4, 39-43.

Kessler, W.R.; Popovic, M.K.; Robinson, C.W. Xanthan Production in an Extemal- Circulation-Loop Airlift Fermenter. The Can. J. Chem. Eng., 1993, 71, 101-106.

Kierstan, M.P.J.; Coughlan, M.P. Immobilization of cells and enzymes by gel entrapment, in "Immobilized Cells and Enzymes: A Practical Approach". Woodward, J. (Ed.), 1985. IRL Press, Washington D.C. pp 39-48.

Kimura, S. Transport phenomena in membrane separation processes. J. Chem. Eng. Japan, 1992, 25(5), 469-479.

Klimek, J.; Ollis, D.F. Extracellular microbial polysaccharides: kinetics of Pseudomonas sp., Azotobacter vinelandii, and Aureobasidium pullulans batch fermentations. Biotechnol. Bioeng. 1980 ,22(11), 2321-2342.

Knapp, J.S.; Howell, J.A. Treatment of primary sewage sludge with enzymes. Top Enzyme Ferment. Biotechnol. 1980, 4, 85.

Kolot, F.B. Immobilized Microbial Systems: Principles, Techniques, and Industrial Applications. 1988. Robert E. Krieger Publishing, Malabar, FL.

Kulkarni, S.S.; Funk, E.W.; Li, N.N. Ultrafiltration. In Membrane Handbook. Ho, W.S.W.; Sirkar, K.K. (Eds.), 1992. Van Nostrand Rcinhold, New York, NY.

Kuu, W.Y.; Polack, J.A. Improving immobilized biocatalysts by gel phase polymerization. Biotechnol. Bioeng, 1983, 25, 1995.

Lasky, M.; Grand, D. Use of microporous hollow fiber membranes in cell harvesting. American Biotechnology Laboratory, 1985, Nov./Dec.

Lehmann, J.K. In Biotechnology. Brauer, H. (Ed.), 1985, Weinheim, FRG: VCH. Vol. 2, p. 607. 210 Leslie, V.L.; Rose, J.B.; Rudkin, G.O.; Fcllzin, J, In New Industrial Polymers. Deanin, R.D. (Ed.), 1974. ACS Symp. Ser. No. 4, Am. Chem. Society, Washington D.C.

Lewis, V.P.; Yang S.-T. Continuous propionic acid fermentation by immobilized Propionibacterium acidipropionici in a novel packed*bed bioreactor. Biotechnol. Bioeng. 1992, 40, 465-474.

Li, B.; Song, L.; Chang, S.; Yang, W. Mass transfer characteristics of the downflow liquid jet loop reactor. Gaoxiao Huaxue Gongcheng Xuebao , 1991, 5(2), I '>2-7.

Lilly V.G.; Wilson, H.A.; Leach, J.G. Bacterial polysaccharid , (II) Lab.-scale production of polysaccharides by species of Xanthomonas. Appl. Microbiol. 1958, 6, 105-108.

Lim, T.; Uhl, J.T.; Prud'homme, R.K. Rheology of self-associated concentrated xanthan solutions. J. Rheology, 1984, 28(4), 367-379.

Lloyd, D.R.; Meluch, T.B. Selection and evaluation of membrane materials for liquid separations. In Material Science of Synthetic Membranes. Lloyd, D.R. (Ed.), 1985. ACS Symp. Ser. No. 269, Am. Chem. Society, Washington D.C.

Lo, Y.-M. Effects of medium composition on product formation and cell growth in xanthan gum fermentation. Master's Thesis, 1993.

Long, R.; Roubicck, R. Centrifugal film reactor - a new concept for cell cultivation. Bioprocess Eng. 1988, 3, 73-77,

Mandelkem, L.; Flory, P.J. Dependence of the diffusion coefficient on concentration in diluted polymer solutions. J. Chem. Phys. 1952, 20, 212-215.

Margaritis, A.; Pace, G.W. Microbial polysaccharides. In Comprehensive Biotechnology, vol. 3; Moo-Young, M. (Ed.), 1985; Pergamon Press, New York, NY; pp 1005-1043.

Margaritis, A.; Wilke, C.R. Study of the rotorfermentor and the kinetics of ethanol fermentation. Biotechnol. Bioeng. 1978, 20, 709.

Margaritis, A,; Zajic, J.E. Mixing, mass transfer, and scale-up of polysaccharide fermentations. Biotechnol. Bioeng. 1978, 20(7), 939-1001.

Maury, L.G. Production of xanthan gum by emulsion fermentation. U.S. Patent 4,352,882. October 5, 1982.

McLean, R.J.C.; Caldwell, D.E.; Costerton, W. Biofilms, Naturally Occurring Communities of Immobilized Cells. In Immobilized Biosystems: Theory and Practical Applications. Veliky, I.A.; McLean, R.J.C. (Eds.), 1994. Blackie Academic & Professional, Glasgrow, UK.

McNeely, W.H.; Kang, K.S. In Industrial Gums. Whistler, R.L. (Ed.), 1973; Academic Press, New York, NY; pp 476-497.

McNeil, B.; Harvey, L.M. Viscous Fermentation products. Critical Reviews in Biotechnol., 1993, 13(4), 275-304. 211 Meyer, E.L.; Fuller, G.G.; Clark, R.C.; Kulicke, W.M, Investigation of Xanthan Gum Solution Behavior under Shear Flow Using Rheooptical Techniques. Macromolecules , 1993,26, 504-511.

Milas, M.; Rinaudo, M.; Tinland, B. The viscosity dependence on concentration, molecular weight and shear rate of xanthan solutions. Polymer Bulletin 1985, 14, 157- 164.

Misra, T.K.; Bamett, S.M. Evaluation of a novel foam fermentor in the production of xanthan gum. Biotechnol. Progresses, 1987, 227-237.

Mohr, R.J.; Khan, A.S. Higee - a New Approach in Groundwater Clean-up. Proceedings of the AQTE Conference, Montreal, Canada, 1987.

Moraine, R. A.; Rogovin, P. Kinetics of the xanthan fermentation. Biotechnol. Bioeng. 1 9 7 3 , 15, 225-237.

Moraine, R. A.; Rogovin, P. Xanthan biopolymer production at increased concentration by pH control. Biotechnol. Bioeng. 1971, 13, 381-391.

Morris, E.R.; Rees, D.A.; Young, G.; Walkinshaw, M.D.; Darke, A. Order-disorder transition for a bacterial polysaccharide in solution: a role for polysaccharide conformation in recognition between Xanthomonas pathogen and its plant host. Journal of Molecular Biology. 1 9 7 7 , 110(1): 1-16.

Mosbach, K.H.; Birnbaum, S.; Hardy, K.; Davies, J.; Buelow, L. Formation of proinsulin by immobilized Bacillus subtilis. Nature (London), 1983, 302(5908), 543.

Moser, A.; Mayr, B.; Jury, W.; Steiner, W.; Horvat, P. Mathematical models for mixing in deep jet bioreactors: analysis. Bioprocess Engineering, 1991, Vol. 7.

Mudgett, R.E. Controlled gas environments in industrial fermentations. Enzyme Microh. Technol., 1980, 2, 273-280.

Mukhopadhyay, S.N.; Das, D.K. In Oxygen Responses, Reactivities, and Measurements in Biosystems. 1994; CRC Press, Ann Arbor, MI; pp 127-178.

Mulder, M. In Basic Principles o f Membrane Technology, Ch. 6 . Kluwer Academic Publishers, 1991, pp 198-212.

Nakajima, S.; Funahashi, H.; Yoshida, T. Xanthan gum production in a fermentor with twin impellers. J. Ferment. Bioeng. 1990. 70, 392-397.

Ni, X.; Gao, S.; Pritchard, D.W. A study of mass transfer in yeast in a pulsed baffled bioreactor. Biotechnol. Bioeng. 1995 ,45, 165-175.

Nienow, A.W. Mixing studies on high viscosity fermentation processes xanthan gums. The World Biotech Report , 1984, 1, 293-304.

Norton, C.J.; Falk, D.O.; Luetzelschwab, W.E. Xanthan biopolymer semipilot fermentation. Soc. Pet. Engineers J., 1981, 21, 205-217. 212 Norton, I.T.; Goodall, D.M.; Frangon, S.A.; Morris, E.R.; Rees, D.A. Mechanism and dynamics of conformational ordering in xanthan polysaccharide, J. Mol. Biol. 1984, 175, 371-394.

Oda, G.; Samejima, H.; Yamada, T. Proceedings of 'Biotech. 83\ Online Publ., Northwood, p597.

Ostwald, W. Z Phys. Chem. 1 9 2 4 ,4 7 //, 62.

Ostwald, W.Z.; Auerbach, R. Kolloid-Zeit. 1926, 38, 361.

Oviatt Jr., H.W.; Brant, D.A. Viscoelastic behavior of thermally treated aqueous xanthan solutions in the semidilute concentration regime. Macromolecules, 1994,27, 2402-2408.

Pace, G.W.; Righelato, R.C. Production of extracellular microbial polysaccharides. In Advances in Biochemical Engineering, vol. 75; Fiechter, A. (Ed.), 1980; Springer- Verlag: New York; pp 41-70.

Parts, A.G.; Elbing, E. The polymerisation of acrylonitrile in centrifugal fields. Paper presented at American Chemical Society Meeting, Philadelphia. Polym. Prepr. 1975, 16, 211-216.

Patton, J.T. U.S. Patent 3,964,972; 1976.

Peters, H.-U.; Herbst, H.; Hesselink, P.G.M.; Lunsdorf, H.; Schumpe, A.; Deckwer, W.-D. The influence of agitation rate on xanthan production by Xanthomonas campestris. Biotechnol. Bioeng. 1989 ,34, 1393-1397.

Peters, H.-U.; Suh, I.-S.; Schumpe, A.; Deckwer, W.-D. Modeling of batchwise xanthan production. The Canadian Journal o f Chemical Engineering, 1992, 70, 742-750.

Phillips, C.R.; Poon, Y.C. Immobilization of cells. In Biotechnology Monograph, Vol. 5. Aiba, S.; Fan, L.T.; Fiechter, A.; Klein, J.; Schiigerl, K. (Eds.), 1988. Springer-Verlag, New York, NY.

Pickering, K.D.; Wiesner, M.R, Cost model for low-pressure membrane filtration. J. Environ. Eng., 1993, 119(5), 772-797.

Pinches, A.; Patient, L.J. Rate and yield relationships in the production of xanthan gum by batch fermentation using complex and chemically defined media. Biotechnol. Bioeng. 1986 ,2 8 , 1484-1496.

Pons, A.; Dussap, C.G.; Gros, J.B. Modelling Xanthomonas campestris batch fermentation in a bubble column. Biotechnol. Bioeng. 1989, 33, 394-405.

Qadeer, M.A.; Baig, S. Microbial synthesis of xanthan gum by repeated-batch process. Pakistan Journal of Scientific and Industrial Research, 1987, 30(12), 886-889.

Rinaudo, M.; Milas, M. Polyelectrolyte behavior of a bacterial polysaccharide from Xanthomonas campestris: comparison with carboxymethyl-cellulose. Biopolymers, 1978, Vol. 17. 213 Robinson, D.K.; Wang, D.I.C. A novel bioreactor system for biopolymer production. Annals N. Y. Acad. Sci. 1988,506, 229-241.

Robinson, D.K.; Wang, D.I.C. A transport controlled bioreactor for the simultaneous production and concentration of xanthan gum. Biotechnol. Prog. 1988, 4, 231-241.

Rochefort, W.E.; Middleman, S. Rheology of xanthan gum: salt, temperature, and strain effects in oscillatory and steady shear experiments. J. Rheology, 1987, 31(4), 337-369.

Rogovin, S.P.; Albrecht, W.J.; Sohns, V. Production of industrial grade polysaccharide B-1459. Biotechnol. Bioeng. 1965,7, 161-169.

Rogovin, S.P.; Anderson, R.F.; Cadmus, M.C. Production of polysaccharide with Xanthomonas campestris. J. Biochem. Microbiol. Technol. Eng. 1961, 3, 51.

Roman, R.V.; Gavrilescu, M. Oxygen transfer efficiency in the biosynthesis of antibiotics in bioreactors with a modified Rushton turbine agitator. Acta Biotechnol. 1994, 14(2), 181-192.

Roubicek, R.; Feres, V. Equipment for enhanced mass transfer in chemical and biochemical processes. U.S. Patent 4,657,677; 1987.

Sanderson, G.R. Applications of xanthan gum. Brit. Polym. J. 1981, 13, 71-75.

Scheraga, H.A.; Mandelkem, L. Hydrodynamic properties of proteins. J. Am. Chem. Soc. 1953, 75, 179-184.

Schiigerl, K. Neue bioreaktoren fUr aerobe prozesse. Chem. Ing. Tech. 1980, 52, 951- 965.

Schumpe, A.; Diedrichs, S.; Hesselink, P.G.M.; Nene, S.; Deckwer, W.D. Xanthan production in emulsions. Biochemical Engineering - Stuttgart, (Proceedings. International Symposium on), 2nd, Meeting Date 1990, 196-9, 1991.

Schweickart, R.W.; Quinlan, A.V, Kinetics of xanthan production when NH 3 -N limits biomass synthesis and glucose limits polysaccharide synthesis. J. Biomechanical Eng., 1989, 111, 166-172.

Setford, S.J.; Barker, P.E.; Ganetsos, G. Combined bioreaction and separation in centrifugal fields. Bioseparation 1994,4, 117-126.

Shiloach, J.; Kaufman, J.B. Hollow fiber microfiltration methods for recovery of rate Basophilic Leukemia cells from tissue culture media. Biotechnol. Prog. 1986 ,2 (4), 230- 233.

Shu, C.H. Multiphase bioreactors for recombinant yeast fermentation. Ph.D. dissertation. The Ohio State University, Columbus, Ohio, 1992.

Shu, C.H.; Yang, S.T. Effect of temperature on cell growth and xanthan production in batch cultures of Xanthomonas campestris. Biotechnol. Bioeng. 1990, 35, 454-468. 214 Shu, C.H.; Yang, S.T. Kinetics and modeling of temperature effects on batch xanthan gum fermentation. Biotechnol. Bioeng. 1991, 37, 567-574.

Shular, M.L.; Cho, T.; Wickham, T.; Ogonah, O.; Kool, M.; Hammer, D.A.; Granados, R.R.; Wood, H.A. Bioreactor development for production of viral pestcides or heterologous proteins in insect cell cultures. Ann. N.Y, Acad. Sci. 1990, 589, 399-422.

Shuler, M.L.; Kargi, F, In Bioprocess Engineering: Basic Concepts. 1992. Prentice Hall, Englewood Cliffs, NJ. pp 340-347,

Silman, R.W.; Rogovin, P. Continuous fermentation to produce xanthan biopolymer: Laboratory investigation. Biotechnol. Bioeng. 1970, 12(1), 75-84.

Sittig, W. Fermentation reactors. CHEM TECH, 1983, 13(10), 606-613.

Sloneker, J.H.; Jeanes, A.R. Exocellular bacterial polysaccharide from Xanthomonas campestris NRRL B-1459: (I) Constitution. Can. J. Chem, 1962, 40, 2066-2071.

Smelser, S.C.; Woodcock, K.E.; Meyer, H.S.; Fowler, R. Selective acid gas removal using the Higee absorber. Proceedings of AIChE Meeting, Orlando, FL, 1990.

Smith, I.H.; Pace, G.W. Recovery of microbial polysaccharides. J. Chem. Technol. Biotechnol. 1982 ,3 2 , 119.

Smith, I.H.; Symes, K.C.; Lawson, C.J.; Morris, E.R. The effect of pyruvate on xanthan solution properties. Carbohydr. Polym. 1984 ,4(2), 153-157.

Solomon, J.; Elson, T.P.; Nienow, A.W. Cavern sizes in agitated fluids with a yield stress. Chem. Eng. Commun. 1981, 11, 143-164.

Sonsbeek, H. M., Tuin, S. P.; Tramper, S. Mixing in liquid-impelled loop reactors. J. Biotechnol. Bioeng. 1992, 39.

Souw, P.; Demain, A.L. Nutritional studies on xanthan production by Xanthomonas campestris NRRL B-1459. Appl. Environ. Microbiol, 1979, 37, 1186-1192.

Stanbury, P.F.; Whitaker, A. Aeration and agitation. In Principles of Fermentation Technology. Pregamon Press: New York; 1984. Chapter 9, pp 169-192.

Stankowski, J.D.; Mueller, B.E.; Zeller, S.G. Location of a second O-acetyl group in xanthan gum by the reductive-cleavage method. Carbohydrate Res., 1993, 241, 321-326.

Stauffer, K.R.; Leeder, J.G. Extracellular microbial polysaccharide production by fermentation on whey or hydrolyzed whey. J. Food Sci. 1978, 43, 756.

Stokke, B.T., Elgsaeter, A.; Smidsrold, O. Electron microscopic study of single- and double-stranded xanthan. Int. J. Biol. Macromol.1986, 8, 217-225.

Suh, I.S.; Herbst, H.; Schumpe, A.; Deckwer, W.-D. The molecular weight of xanthan polysaccharide produced under oxygen limitation. Biotechnology Letters, 1990, 12(3), 201-206. 215 Suh, I.S.; Schumpe, A.; Deckwer, W.D. Xanthan production in bubble column and air-lift reactors, Biotechnol. Bioeng . 1991, 39, 85-94.

Sutherland, I.W. and Ellwood, D.C. Microbial exopoly saccharides: industrial polymers of current and future potential. In Microbial Technology: Current State, Future Prospects. Cambridge U. Press, 1979.

Sutherland, I.W. Biosynthesis of microbial exopolysaccharide. In Advances in Microbial Physiology. A. H. Rose (Ed.), 1982, Academic Press, 79-150.

Sutherland, I.W. Extracellular polysaccharides. In Biotechnology, vol. 3; 1979, Chapter 3h. Academic Press, New York, NY.

Sutherland, I.W, Structure-function relationships in microbial exopolysaccharides. Biotech. Adv., 1994, 12, 393-448.

Sutherland, I.W. Surface Carbohydrates of The Prokaryotic Cell. Academic Press, New York, NY. 1977.

Svensson, H.; Hagdahl, L.; Lemer, K.D. Zone electrophoresis in a density gradient. Stability conditions and separation of serum proteins. Science Tools , 1957 ,4 , 1-10.

Swift, C.J.; Atkins, W.S. Alcohol recovery by distillation. Monograph: European Brewery Convention, 1983, 185-202.

Symes, K.C. The relationship between the covalent structure of the Xanthomonas polysaccharide (xanthan) and its function as a thickening, suspending, and gelling agent. Food Chem. 1980, 6(1), 63-76.

Tait, M.I.; Sutherland, I.W.; Clarke-Sturman, A.J. Effect of growth conditions on the production, composition and viscosity of Xanthomonas campestris exopolysaccharide. J. Gen. Microb. 1986, 132, 1483-1492.

Tam, K.C.; Tiu, C. Improved correlation for shear-dependent viscosity of polyelectrolyte solutions. J. Non-Newtonian Fluid Mechanics, 1993, 46, 275-288.

Tam, K.C.; Tiu, C. Rheology of water-soluble polymers: a comparative study on the effect of monovalent salt. Polym.-Plast. Technol. Eng., 1993 ,32(1&2), 123-138.

Taylor, K.C.; Nasr-El-Din, H.A. Xanthan biopolymers: a review of methods for the determination of concentration and for the measurement of acetate and pyruvate content J. Pet. Sci. Eng., 1993, 9, 273-279.

Tecante, A.; Choplin, L. Gas-liquid mass transfer in non-Newtonian fluids in a tank stirred with a helical ribbon screw impeller. The Can. J. Chem. Eng., 1993, 71, 859-865.

Thonart, Ph.; Paquot, M.; Hermans, L.; Alaoui, H. Xanthan production by Xanthomonas campestris NRRL B-1459 and interfacial approach by zeta potential measurement. Enzyme Microb. Technol. 1985, 7, 235-238. 216 Thome, L.; Tansey, L.; Pollock, TJ. Direct utilization of lactose in clarified cheese whey for xanthan gum synthesis by Xanthomonas campestris. J. Ind. Microbiol. 1988, 3(5), 321-328.

Torres, L.G.; Brito, E.; Galindo, E.; Choplin, L. Viscous behaviour of xanthan aqueous solutions from a variant strain of Xanthomonas campestris. J. Ferm. Bioeng., 1993, 75(1), 58-64.

Tweddle, T. A.; Striez, C.; Tam, C. M.; Hazlett, J. D. Polysulfone membranes I. Performance comparison of commercially available ultrafiltrarion membranes. Desalination, 1992 ,8 6 , 27-41. van Oers, C. W.; Vorslman, M. A. G.; Muijselaar, W. G. H. M.; Kerkhof, P. J. A. M. Unsteady-state flux behavior in relation to the presence of a gel layer. J. Membrane Sci. 1992, 73, 231-246.

Van't Riet, K. Review of measuring methods and results in non-viscous gas-liquid mass transfer in stirred vessels. Ind. Eng. Chem. Process Des. Dev. 1979, 18(3), 357-360.

Vashitz, O.; Sheintuch, M. Analysis of polymer synthesis rates during steady-state growth of X. campestris. Biotechnol. Bioeng. 1991, 37, 383-385.

Vickroy, T.B., Blanch, H.W.; Wilke, C.R. Growth of prokaryotic cells in hollow-fibcr reactors. Biotechnol. Letters, 1982, 4, 483.

Viehweg, H. Process for the preparation of exocellular biopolymers. U.S. Patent 4,871,665. October 3, 1989.

Wada, M.; Kato, J.; Chibata, I. A new immobilization of microbial cells. Immobilized growing cells using carrageenan gel and their properties. J. Ferment. Technol. 1980, 58, 327.

Wakeman, R.J.; Tarleton, E.S. Membrane fouling prevention in crossflow microfiltration by the use of electric field. Chem. Eng. Sci. 1987, 42(4), 829-942.

Wang, S.-S. Effect of solution viscosity on ultrafiltration flux. J. Membrane Sci. 1988, 39, 187-194.

Wang, D.I.C., Cooney, C.L., Demain, A.L., Dunnill, P., Humphrey, A.E., Lilly, M.D, 1979. Fermentation kinetics, pp. 57-97. In C.G. Heddn (ed.), Fermentation and enzyme technology. John Wiley & Sons, New York.

Weiss, R.M.; Ollis, D.F. Extracellular microbial polysaccharides. I. Substrate, biomass, and product kinetic equations for batch xanthan gum fermentation. Biotechnol. Bioeng. 1980, 22, 859-873.

Whitcomb, P.J.; Macosko, C.W. Rheology of xanthan gum. J. Rheol. 1978, 22(5), 493- 505.

Whitfield, C.; Sutherland, I.W.; Cripps, R.E. Glucose metabolism in Xanthomonas campestris. J. Gen. Microbiol. 1981, 124, 385-392. 217 Wilkins, M.J.; Davies, M.C.; Jackson, D.E.; Mitchell, J.R.; Roberts, C.J.; Stokke, B.T.; Tendter, S.J.B. Comparison of scanning tunnelling microscopy and transmission electron microscopy image data of a microbial polysaccharide. Ultramicroscopy , 1993, 48, 197- 201.

Yang, S.T.; Guo, M. A Kinetic model for methanogenesis form whey permeate in a packed bed immobilized cell bioreactor. BiotechnoL Bioeng . 1991, 37, 375-382.

Yang, S.T.; Zhu, H.; Li, Y. Continuous propionate production from whey permeate using a novel fibrous bed bioreactor. Biotechnol. Bioeng. 1994, 43(11), 1124-1130.

Yeh, H.-M.; Cheng, T.-W. Osmotic-pressure model with permeability analysis for ultrafiltration in hollow-fiber membrane modules. Separation Technol . 1993, J, 91-98.

Zaidi, A.; Ghosh, P.; Schumpe, A.; Deckwer, W.-D. Xanthan production in a plunging jet reactor. Appl. Microbiol. Biotechnol. 1991, 35, 330-333.

Zatz, J.L. Applications of gums in pharmaceutical and cosmetic suspensions. Ind. Eng. Chem. Prod. Res. Dev., 1984, 23(1), 12-16.

Zhu, H. A novel fibrous-bed bioreactor for mammalian cell culture. Ph.D. dissertation, The Ohio State University, Columbus, Ohio, 1995. APPENDIX A

FERMENTA TION ANALYTICAL METHODS

AJ SAMPLE ANALYSIS PROCEDURES

Sample analysis procedures are illustrated in a flow diagram (Figure A. 1). After centrifuging, the cell-free supernatant was stored in a freezer until analysis for glucose, viscosity, and Final product concentration. The cell precipitate was resuspended in water and analyzed immediately for cell density determination. Xanthan gum was isolated from the cell-free supernatant by alcohol precipitation.

A.2 CELL DENSITY DETERMINATION

Depending on the broth viscosity, the broth samples were diluted with tap water by a factor of two to six, and the diluted solutions were centrifuged using Beckman Model J2-

21 Centrifuge at 12,000 rpm (16,000 x g) for 30 minutes at 5°C to precipitate the suspended cells. Cells were then resuspended in water and the optical density at 650 nm was measured using a spectrophotometer (Sequoia-Tumer Model 340). The readings were referred back to a predetermined calibration curve (Figure A.2) to estimate the cell density.

218 219

Samples

Centrifuge

Prod pit ales Supernatants

Cell density Viscosity Glucose concentration Precipitation determination determination determination by Ethanol

I Spectrophotometer 1 Brookfield Glucose Analyzer Centrifuge I at 650 nm J Viscometer YS1 2700

f "\ Cell viability (Plate count) V J Xanthan concentrationj determination I

FIGURE A.I: Sample analysis procedures. FIGURE A.2: FIGURE

Cell Density (g/L) Calibration curves for cell density and plate count numbers versus OD. versus numbers count plate and density cell for curves Calibration 0.00 0.05 0.10 pf'»>i |in»TMi i n i m 0,20 i nrr

220 221

The calibration curve for cell density determination was obtained as follows. Cells

obtained from previous fermentations were collected from 250 ml broth by centrifugation.

The collected cells were then resuspended in a small amount of water, ca. 10 ml. A series

of dilution were then made from the concentrated cell suspension. The total dry weight of

cells in the concentrated solution was also determined after drying at 107*^2 for about 7

hours. This procedure was repeated to reduce the experimental errors to within 0.2%. A

linear relationship between the optical density and cell density was then obtained with the

optical density less than 0.5.

The correlation between colony numbers counted on the growth medium agar plate

from free cells of Xanthomonas campestris and the corresponding optical density (at 650

nm) is also given in Figure A.2. Technical information regarding how this correlation was

used in determining the relative cell viability of the immobilized cells of X. campestris has

been described in detail in Chapter III.

A.3 GLUCOSE ANALYSIS

The glucose consumption rates in fermentations were determined by measuring the glucose concentration in samples. Glucose concentrations in cell-free supernatant were determined by the YSI glucose analyzer (Model 2700 SELECT) in the Fermentation

Laboratory at Microbiology Department of the Ohio State University. As a single channel unit, the 2700 SELECT provides 90-second measurements of glucose (detection range; 0-

25 g/1) simply by presenting properly diluted cell-free samples to the needle for automatic sample aspiration (adjustable from 5 to 65 microliters). 222 This method is quite reliable and has become a routine procedure not only in

microbiological laboratories but also in hospitals. The analysis is based on a biosensor

membrane; a three-layer laminate with immobilized glucose oxidase in the middle layer.

The probe, covered by the membrane, is situated in a sample chamber where a flow of

buffer solution containing the glucose sample is pumped past the sensor surface. Some of

the glucose diffuses through the membrane. When glucose in the sample contacts the immobilized glucose oxidase, it is rapidly oxidized, producing hydrogen peroxide. The

hydrogen peroxide is, in turn, oxidized at the platinum anode, producing electrons. The electron flow is linearly proportional to the local H 2O2 concentration and, therefore, to the glucose concentration. Measurements are virtually unaffected by sample color, turbidity, density, viscosity, pH, volatility, specific gravity, temperature, index of refraction, optical activity, or the presence of other nutrients. The enzymes can be reused for up to a month since they are immobilized. There are some noices affecting the baseline calibration while samples containing over-range glucose conentrations are aspirated.

A.4 XANTHAN CONCENTRATION

Xanthan concentration in the fermentation broth was estimated from the broth viscosity according to the linearity between xanthan concentration and broth viscosity at low concentration. After ceil removal from the broth, the solution viscosity was correlated with the amount of xanthan gum present in the solution. About 30 ml of cell-free supernatant from the broth sample was diluted by 6-7 times, and the solution viscosity was measured on a Brookfield viscometer (RVTD II) with spindle No. 1 at 100 rpm. The concentration of xanthan gum was then estimated in comparison with a calibration curve

(Figure A.3), which was obtained from standard xanthan solutions. Viscometer Readings (Spindle #1, 100rpm) FIGURE A.3: FIGURE : 8.1556 + 14.504X 14.504X + 8.1556 : 1 HimIII I| HH m I11 i III III H IT II tl l| III IM I III [IIM H ata u ocnrto determination. concentration gum Xanthan R*2 R*2 0.990 = 223 224

Xanthan concentration was also determined by measuring the dry weight of

xanthan gum isolated and purified from the fermentation broth by alcohol precipitation.

About 20-30 ml of cell-free supernatant from the broth sample was precipitated with 2-3

volumes of alcohol, and saturated KCI solution (ca. 5% of total volume) was added to

precipitate xanthan gum. After centrifugation, the precipitated gum was dried in an oven,

and then weighted to estimate xanthan concentration in the broth.

A.5 FINAL CELL AMOUNT IN CPBR

Bioreactor CPBR-LC (5 liters) CPBR-GC (2.5 liters)

Sample # 1 2 3 4 5 6

OD (8 x) 0.17 0.16 0 .2 0 0.24 0.26 0.30

Cell (g/L) 0.070 0.068 0.076 0.095 0 .1 0 2 0.121

Water (L) 0 .2 0.2 0.2 0 .2 0 .2 0 .2

Cell (g) 0 .1 1 2 0.109 0.122 0.152 0.163 0.194

Area(%) 0.33 0.32 0.36 0.42 0.41 0.50

Total (g) 34.0 38.4 225

A.6 ULTRAFILTRATION MEMBRANE AREA CALCULATION

at 70 psig to process 1 L xanthan = 25 g xanthan # 1 # 2 $400 $200 Shear (1/s) Q(m/min) time (s) kg/year Mem$/kg/yr Mem$/kg/yr 4.90E+02 1011.6 12191 17.71798 1.46743 0.73371 3.92E+02 809.1 13095 16.49484 1.57625 0.78812 3.18E+02 656.6 14000 15.42857 1.68518 0.84259 1.87E + 02 386.4 1 6568 13.02146 1.99670 0.99835 1.53E+02 315.3 17704 12.20063 2.13103 1.06551 1.24E+02 255.0 18946 11.39962 2.28077 1.14038 9.91 E+01 204.5 20334 10.62260 2.44761 1.22380 7.94E+01 163.8 21829 9.89509 2.62756 1.31378 5.94E + 01 122.5 23958 9.01577 2.88383 1.44191

Electricity $0.05/KWh Power (W) P$/kg #1 Total $ #2 Total $ 209.42 1.418355 2.88579 2.152072 135.16 0.983289 2.55953 1.771414 107.18 0.833622 2.51880 1.676214 59.48 0.548141 2.54484 1.546493 47.45 0.466697 2.59773 1.532215 37.49 0.394644 2.67542 1.535033 29.34 0.331444 2.77905 1.555249 22.94 0.278198 2.90576 1.591980 16.61 0.221079 3.10491 1.662995 APPEN D IX B

EQUATION DERIVATION

B.l PRESSURE DROP IN ULTRAFILTRATION MODULE

First, set-up and solve the differential equations for the pipe flow of a non-

Newtonian fluid which is described by power law:

dU , | (B.l) where K and n are constants. The flow is isothermal and fully developed, and the fluid is incompressible.

Steady state pipe flow (no external forces)

(1) One dimenlional flow: U, = Ue = 0 (B.2)

(2) Symmetry: — = 0 (B.3) dQ

dU (3) Incompressible: V ■ U - 0 => — *- = 0 => U - f{r) (B.4) dz 1

226 227

Then, the momentum equations in terms of rfor this system become:

dP . r-component: t _ 0 (B.5)

z-component: (B.6 ) Td r “ ( r r ) d z

- i - m

dP, _ I d (B.7) dz r dr -i-fj dPt _ K d (B.8 ) dz dr (-£)'

Boundary conditions: r = rn, U = 0 (B.9)

r = 0, ^ = 0 (B.10) dr

2 = 0 , Pt = Pm (B .ll)

z = L . P = Pol + A/>, (B.12) where z = 0 is within the fully developed flow region. Integrating Equation (B. 8 ) over z yields

• L *****' dP,Jn = ——K d dz (B.l 3) -I r dr

d_ r A Pt (B. 14) dr t - f l ' K L 228 Integrating Equation (B.14) over r yields

2 AP. L + C, (B.15) \ dr ) 2i2 K L from the second B.C., Cj ~ 0

(B.l 6 ) dr \ 2K L J { 2KLJ where

(B.17) =l f-^ 2KLJ l fT + ] -V T- A -2— r ” +C2 (B .18) n + 1 2 from the first B.C.:

fl+1 n C2~~ A (B.19) n + 1

rt+1 fl+I " n + \ / > 'r_y n r n \ U= A 1 - l - = Uz, m ax (B.20) n + 1 Jo) , ro> where

H + l / A FI \ “ tt+1 U -A n r~- n AM " r ^ ~ (B.21) U l.msa, n + 1 | 0 n + lv ,1 2KLJt 0

The value of Uz, ^ can thus be obtained:

/. = U - 1! n + l n f V r~ (B.22) a r 2 J,ave 1 + 3n 1 + 3fl I 2 KLJ 0 229 Rearrange Equation (B.22) gives

/~ , - 3fl+lV 3,1+1 /I AP, = -2 K L Qr0 " (B.23) v j

Assumptions made:

( 1) laminar flow ( 2 ) steady state

(3) no external forces (i.e., gravity) (4) no entrance effect

(5) incompressible fluid ( 6 ) K and n are constants

B.2 SHEAR RATE AT WALL, VISCOSITY, AND REYNOLDS NUMBER

B.2.1 Shear Rate at Wall

The shear rate at the wall of ultrafiltration membrane tube can be computed from the

Hagen-Poiseuille law for laminar pipe flow;

f f O = ^ = ^ (B.24) \ d r )** n To ro where UL aye is the average velocity across the pipe, and

u *'*" = %S = - r 7tr T 230 B.2.2 Viscosity

Again, from the Hagen-Poiseuille law we can correlate viscosity with shear stress:

(B.26)

By measuring the values of the diameter and the length of the tube, d0 and L, respectively, and the pressure drop of the tube, -APh the viscosity of the fluid in pipe, p , can he determined with the aid of the known value of shear rate at wall.

B.2.3 Reynolds Number, N Re

Specifically, the Reynolds number, NRe, for pipe flow was:

At _ ^ 1,aye P Re ~ (B.27) where p is the fluid density. In this study, the value of NRe was found to locate in the region of laminar flow.

B.3 SHEAR RATE AT WALL IN BROOKFIELD VISCOMETER

The viscosity of a fluid was measured by placing the spindle into a cylindrical container containing the fluid. The spindle was rotated at a constant speed, co , which was low enough that the flow was laminar. A torque was measured at the spindle.

The flow in this measurement was in the 9 direction only, so that 231 £/, =

^- = 0 (B.29) d t

The viscometer was long enough in the e-direction so that end effents were negligible; thus

^ ■ = 0 (B.30)

The flow was symmetric:

fill n i = a

The viscometer configuration was such that gravity acted only in the e-direction, so that

&=0 = ge (B.32)

The value of gz was 9.8 m s 2.

The Navier-Stokes equations for cylindrical coordinates for the r component became: 1 -vl dp (B.33) r p dr

For the 6 component was:

d2U„.ldUa ue 0 = ^ + 7 l f ' ^ (BM) 232 For the z component was:

0 = - - ^ + «, (B.35) p az

Equation (B.35) shows that the vertical pressure drop was due only to gravity acting on the fluid mass:

& = p g , (B.36) dz

Equation (B.34) did not involve a convective contributaiton. It was an exact differential and could be written as

d_ f \d(rUe) ' = 0 (B.37) dr r dr where the partial derivatives were replaced by total derivatives, since all variation here was in a single direction r. After integration, Equation (B.37) became

}_£[H A = c (B.38> r dr

Integrating again:

r 2 rUe = C( — + C3 (B.39)

The two boundary conditions were

u e (r = O = 0 y e (r = ri) = °>r, (B.40) where o> was the velocity of rotation of the spindle. Substitution of the boundsary conditions into Equation (B.39) resulted in two equations in two unknowns:

0 = ci^- + c2 a)rf=Cii- + C2 (B.41) These equations were solved for C/ and C?.

CY = -2(0 r; i {r] - r f) (B.42)

C2^0)(r2or ^ /(r 20~rf) (B.43)

Equations (B.39), (B.42), and (B.43) were combined to give the velocity distribution:

(Or1 (r2 ^ U6=-f~LT ^ -r \ (B.44) ro-nKr J

0>r2 (rl-r1) or (B.45) (rl-rf)

As discussed earlier, the angular velocity co and force on the spindle were measured. The wall shear stress at r = ra was

1 3U A '■H) (B.46) (t/-/r)+7*T

From this study, the term dVT 139 was zero, and Equation (B.46) reduced to

d{U ,lr) U = -H r- (B.47) dr

Combined Equations (B.44) and (B.47) at r = r0 gave

2(0 rj (B.48)

Therefore, the shear rate at wall, yw, was obtained: APPENDIX C

EXPERIMENTAL DATA FOR CHAPTER III

FIGURE 3.2: MATRIX SELECTION

Celt Density (o/L) detected in xanthan fermentation broth Time (hr) 100% C (towel) 100% C (flat) 50% C + 50% PE 100% PE 0 0.00 0.00 0.00 0.00 6 0.11 0.10 0.11 0.10 1 2 0.40 0.35 0.45 0.42 1 7 0.66 0.70 0.68 0.65 23 1.20 1.12 1.23 1.20 26 1.21 1.19 1.20 1.21 36 0.13 0.65 1.10 1.13 43 0.06 0.36 1.00 1.08 50 0.02 0.21 0.90 0.96

234 235

FIGURE 3.4: TYPICAL TWO-STAGE, REPEATED-BATCH

FERMENTATION

Time (hr) Glucose (a/L) Xanthan (a/L) CD (a/L) pH Temp (C DOT (%) rpm 0.0 24.80 0.00 0.00 6.0 23.0 100 1 50 4.0 23.40 0.50 0.10 5.9 23.0 96 1 50 10.0 23.20 1.60 0.30 6.0 23.0 91 150 15.0 23.00 2.50 0.90 6.0 23.0 8 1 150 23.0 19.80 4.00 1.70 6.1 23.0 62 150 24.0 -- - 7.0 30.0 62 350 35.0 12.00 1 1.30 0.20 7.0 30.0 42 350 39.0 8.00 14.80 0.10 7.0 30.0 37 350 43.0 5.05 17.53 0.00 7.0 30.0 25 350 48.0 0.30 21.50 0.00 7.0 30.0 1 5 350 49.0 24.40 0.00 0.00 7.0 30.0 96 350 59.0 19.60 2.50 0.00 7.0 30.0 80 350 65.0 15.60 6.80 0.00 7.0 30.0 63 350 69.5 12.40 8.05 0.00 7.1 30.0 51 350 73.0 1 0.50 12.00 0.00 7.0 30.0 46 350 84.0 3.00 17.53 0.00 7.0 30.0 26 350 86.5 0.02 20.27 0.00 7.0 30.0 1 9 350 236

FIGURE 3.5: ISO RPM FERMENTATION PROFILE

Time (hr) Cell Density (a/L) Glucose (a/L) Xanthan (g/L) 0.0 0.08 25.00 0.00 13.0 0.43 * 0.40 19.0 1.20 20.00 1.60 26.0 1.21 17.70 2.90 49.0 0.13 13.46 8.40 62.0 0.10 11.36 10.60 79.0 0.02 7.60 14.70 93.0 0.00 3.89 18.84 93.0 0.00 30.56 1.80 1 17.0 0.00 25.31 3.60 140.0 0.00 20.65 9.60 165.0 0.00 13.80 14.60 177.0 0.00 9.53 17.50 187.0 0.00 8.06 20.40 187.0 0.00 50.40 2.00 203.0 0.00 44.20 4.50 213.0 0.00 41.70 5.10 225.0 0.00 37.40 7.90 238.0 0.00 33.10 8.60 250.0 0.00 31.00 12.80 262.0 0.00 27.65 15.00 272.0 0.00 23.70 16.10 297.0 0.00 21.15 20.40 333.0 0.00 14.20 24.50 333.0 0.00 59.80 2.40 338.0 0.00 56.60 2.80 345.0 0.00 54.30 4.60 366.0 0.00 50.00 6.20 379.0 0.00 47.70 8.50 392.0 0.00 44.50 1 1.20 405.0 0.00 41.60 12.80 416.0 0.00 38.70 15.70

FIGURE 3.6: ISO RPM FERMENTATION ANALYSIS

Batch Number Yp/s dP/dt 1 0.89 0.20 2 0.83 0.20 3 0.62 0.15 4 0.63 0.16 237

FIGURE 3.7: 350 RPM LIQUID-CONTINUOUS FERMENTATION

Time (hr) Cell Density (g/L) Glucose (g/L) Xanthan (g/L) 0.0 0.00 24.80 0.00 4.0 0.10 23.40 0.50 10.0 0.30 23.20 1.60 15.0 0.90 23.00 2.50 23.0 1.70 19.80 4.00 35.0 0.20 12.00 1 1.30 39.0 0.10 8.00 14.80 43.0 0.00 5.05 17.53 48.0 0.00 0.30 21.50 49.0 0.00 24.40 0.00 59.0 0.00 19.60 2.50 65.0 0.00 15.60 6.80 69.5 0.00 12.40 8.05 73.0 0.00 10.50 12.00 84.0 0.00 3.00 17.53 86.5 0.00 0.02 20.27 87.0 0.00 25.30 0.00 93.0 0.00 21.50 1.60 106.5 0.00 14.80 8.20 112.0 0.00 11.40 10.80 122.0 0.00 3.91 17.53 130.0 0.00 0.95 20.50 131.0 0.00 24.70 0.00 136.0 0.00 20.00 1.60 148.0 0.00 13.60 7.80 154.0 0.00 9.80 12.50 161.0 0.00 6.40 15.25 1 73.0 0.00 0.05 20.80 174.0 0,00 49.50 0.00 181.0 0.00 47.00 5.40 196.0 0.00 35.40 16.53 204.0 0.00 27.00 18.50 212.0 0.00 21.00 26.00 224.0 0.00 9.40 36.00 225.0 0.00 50.30 0.00 234.0 0.00 44.50 5.60 243.0 0.00 36.70 1 1.00 253.0 0.00 27.80 18.20 267.0 0.00 17.90 28.00 276.0 0.00 10.401 35.00 238

FIGURE 3.8: 350 RPM LIQUID-CONTINUOUS ANALYSIS

Batch Number Yp/s dP/dt Viscosity 1 (cP) Viscosity 2 (cP) 1 0.86 0.44 3056 3068 2 0.83 0.54 31 08 3092 3 0.84 0.48 3084 3076 4 0.84 0.50 3072 3058 5 0.90 0.72 31 04 3088 6 0.88 0.69 3096 3090

FIGURE 3.9: 350 RPM GAS-CONTINUOUS FERMENTATION

(See next page) 239

Time (hr) Cell Density (g/L) Glucose (g/L) Xanthan (g/L) 0.0 0.00 25.00 0.00 6.0 0.10 24.60 1.60 14.0 0.80 23.00 2.50 18.0 1.40 21.50 3.60 24.0 1.90 17.50 5.60 36.0 0.13 9.80 12.50 48.0 0.02 1.00 22.50 48.5 0.00 25.10 0.00 52.0 0.00 20.50 4.00 56.0 0.00 15.00 8.00 62.0 0.00 9.50 12.00 66.0 0.00 5.20 17.00 72.0 0.00 0.70 21.50 72.5 0.00 25.10 0.00 76.0 0.00 20.50 4.00 82.0 0.00 15.00 8.00 85.0 0.00 1 1.50 10.50 92.0 0.00 3.80 18.50 96.0 0.00 0.70 21.00 96.5 0.00 25.10 0.00 103.0 0.00 18.50 6.00 109.0 0.00 1 1.50 10.50 1 16.0 0.00 5.20 1 7.00 120.0 0.00 0.70 22.00 120.5 0.00 50.00 0.00 124.0 0.00 46.00 4.50 130.0 0.00 39.50 8.00 136.0 0.00 30.00 15.00 142.0 0.00 23.00 22.00 150.0 0.00 11.20 29.00 152.0 0.00 9.00 32.40 152.5 0.00 50.00 0.00 160.0 0.00 43.00 6.00 164.0 0.00 38.50 9.50 170.0 0.00 29.00 1 7.50 175.0 0.00 22.00 21.50 1 80.0 0.00 15.00 27.00 186.0 0.00 9.00 32.40 186.5 0.00 25.10 0.00 190.0 0.00 19.40 5.50 195.0 0.00 1 1.50 10.50 201.0 0.00 5.20 17.00 205.0 0.00 0.70 20.00 205.5 0.00 25.10 0.00 212.0 0.00 19.50 5.50 218.0 0.00 12.00 1 1.50 222.0 0.00 5.20 17.00 226.0 0.00 0.70 19.50 240

FIGURE 3.10: 350 RPM GAS-CONTINUOUS ANALYSIS

Batch Number Yp/s dP/dt Viscosity 1 (cP) Viscosity 2 (cP) 1 0.94 0.47 3100 3095 2 0.88 0.91 3060 3085 3 0.86 0.89 3050 3075 4 0.90 0.94 3105 3090 5 0.79 1.03 3085 3080 6 0.81 0.97 3080 3070 7 0.82 1.08 3090 3095 8 0.80 0.95 3095 3100

FIGURE 3.12: BATCH FERMENTATION (2.5% GLUCOSE + 0.3% YE)

Time (hr) CD (a/L) Glucose (a/L) Xanthan (a/L) dP/dt/Xs DOT (%) am OUR 0 0.07 25.26 0.52 - 1 00 0.00 0.00 7 0.14 24.99 0.63 0.67 95 -• 1 1 0.25 23.81 1.34 0.48 92 0.12 548.57 20 0.92 20.35 2.27 0.25 88 0.21 225.39 2 5 1.40 17.55 2.85 0.25 87 0.28 200.00 30 1.47 16.71 4.16 0.25 76 -- 34 1.39 15.06 7.89 0.24 69 -- 4 0 1.39 14.31 11.18 0.23 65 0.19 143.33 44 1.34 12.51 12.49 0.23 62 0.18 139.22 50 1.29 9.50 15.73 0.21 60 -- 5 5 1.28 7.44 16.27 0.20 48 0.16 128.00 58 1.30 6.28 18.04 0.17 42 . _ 64 1.21 5.15 19.14 0.12 20 0.09 74.83 6 8 1.29 2.31 20.56 0.09 1 2 -- 72 1.17 2.12 20.98 0.08 4 0.05 38.14 241

FIGURE 3.13: BATCH FERMENTATION (5.0% GLUCOSE + 0.3% YE)

Time (hr) CD (g/U Glucose (g/L) Xanthan (g/L) DOT (%) OTR dP/dt/Xs CUR 0 0.05 49.80 0.29 1 00 0.00 - 0.00 8 0.06 44.40 0.29 95 - 0.66 - 1 4 0.07 43.60 1.17 88 0.06 - 807.82 1 9 0.12 42.40 1.73 84 - 0.76 - 24 0.15 36.70 3.81 80 0.12 - 821.80

30 0.20 36.60 7.91 75 - -- 34 0.35 33.80 9.97 72 0.16 - 460.90 39 0.56 32.80 11.17 61 - 0.52 - 47 0.77 26.82 16.49 56 0.23 0.46 298.67 56 0.98 24.32 19.53 50 0.30 0.36 - 64 1.43 18.94 24.62 44 - 0.27 - 7 1 1.85 17.36 29.31 39 0.20 0.22 107.87 80 1.89 14.40 32.83 20 0.16 0.22 84.88 87 1.86 9.94 36.06 1 0 - 0.22 - 103 1.86 2.51 38.69 8 0.06 0.20 32.26 1 1 8 1.77 0.09 39.57 6 0.04 0.19 22.64

FIGURE 3.14: VOLUMETRIC AND SPECIFIC PRODUCTIVITIES

Volumetric Productivity Specific Productivity Reactor Type 2.5/0.3 5.0/0.3 2.5/0.3 5.0/0.3 STBR 0.30 0.45 0.20 0.24 CPBBR-LC 0.50 0.70 0.07 0.10 CPBBR-GC 0.90 1.00 0.06 0.07 242

FIGURE 3. IS: dP/dt/Xs vs. OUR

STR-DT (2.5% Glu) STR-DT (5.0% Glu) For CPBR fermentation OUR (mg) dP/dt/Xs OUR (mg) dP/dt/Xs OUR (mg) dP/dt/Xs 2.00E+02 2.80E-01 4.61 E+02 5.70E-01 4.56E+01 7.0E-02 1.80E + 02 2.46E -01 3.75E + 02 5.15E-01 5.83E+01 8.2E-02 1.55E+02 2.32E-01 2.99E+02 4.10E-01 7.60E + 01 1 .OE-01 1.43E+02 2.20E-01 2.00E+02 3.10E-01 9.80E+01 1 .2E-01

1.39E + 02 2.1 0E-01 1.45E+02 2 .6 9 E -0 1 --

1.31E + 02 2.00E -01 1.00E + O2 2 .2 3 E -0 1 - -

1.28E+02 1.80E-01 8.49E+01 2 .0 0 E -0 1 - -

1.10E+02 1.65E-01 6.50E+01 1.90E-01 - -

8.00E+01 1.23E-01 3.23E+01 1.85E-01 --

6.00E + 01 9.46E-02 2.26E+01 1.80E-01 -- 4.00E+01 8.13E-02 - -- - APPENDIX D

EXPERIMENTAL DATA FOR CHAPTER IV

FIGURE 4.2: STATIC GASSING-OUT TECHNIQUE

Time (s) DOT (%) Time (s) DOT (%) 0 100.0 313 52.9 1 2 100.0 327 66.0 24 36.4 341 74.0 36 6.4 356 79.0 48 1.8 370 81.3 60 0.0 384 82.9 72 0.0 398 84.4 84 31.4 412 84.4 96 62.7 424 30.7 108 78.2 436 5.4 120 87.7 448 1.5 132 93.6 460 0.0 144 96.4 516 0.0 156 98.2 540 24.1 168 100.0 564 48.3 180 100.0 588 60.2 192 36.4 612 67.6 204 6.4 636 72.1 216 1.8 660 74.2 228 0.0 684 75.6 284 0.0 708 77.0 299 26.5 732 77.0

243 244

FIGURE 4.3: SATURATED DISSOLVED OXYGEN CONCENTRATION

Xanthan (a/L) Saturation % C* fmM/U 0.0 100.0 0.244 12.0 84.7 0.206 22.0 77.3 0.188 35.0 76.0 0.185

FIGURE 4.4(a): kLa IN STR-DT

rpm Water 1.2% Xanthan 2.2% Xanthan 3.5% Xanthan 200 0.0273 --- 400 0.0442 0.0041 0.0050 0.0048 600 0.0556 0.0274 0.01 1 1 0.0054 800 0.0499 0.0443 0.0272 0.0055 1000 0.0483 - - -

FIGURE 4.4(b): kLa IN STR-MP

rpm Water 0.7% Xanthan 1.1% Xanthan 2.4% Xanthan 200 0.0038 - - - 400 0.0057 0.0027 0.0016 0.0021 600 0.0098 0.0038 0.0036 0.0039 800 0.0112 0.0093 0.0078 0.0070

1 000 0.0151 - - -

FIGURE 4.4(c): kLa IN STR-WIO

rpm Water Water/Oif 1.2% Xanthan/Oil 2.2% Xanthan/Oil 3.5% Xanthan/Oil 200 0.027 - - • - 400 0.044 0.046 0.009 0.007 0.006 600 0.056 0.057 0.030 0.020 0.011 800 0.050 0.058 0.050 0.038 0.018

1000 0.048 - - - - 245

FIGURE 4.4(d): kLa IN CPBR-LC

kLa Rotation (rpm) Water 1.2% Xanthan 2.2% Xanthan 3.5% Xanthan 40 0.014 0.008 0.007 0.006 80 0.017 0.012 0.010 0.007 1 20 0.021 0.014 0.013 0.008 150 0.024 0.016 0.016 0.010 200 0.034 0.025 0.018 0.012 250 0.043 0.029 0.022 0.013 300 0.044 0.033 0.026 0.015 350 0.045 0.036 0.030 0.016 400 0.046 0.036 0.031 0.018 450 0.045 0.037 0.030 0.019 500 0.047 0.036 0.031 0.021

FIGURE 4.5: EFFECT OF CIRCULATION IN CPBR

Rotation (rpm) Water (w/) Water (w/o) 1.2% (w/) 1.2% (W/o) 40 0.0150 - 0.0076 - 80 0.0175 0.0100 0.0115 0.0055 1 20 0.0220 - 0.0135 - 150 0.0260 0.0200 0.0165 0.0085 200 0.0360 - 0.0260 - 250 0.0435 0.0400 0.0295 0.0235 300 0.0460 - 0.0350 - 350 0.0480 0.0450 0.0410 0.0380 400 0.0485 - 0.0415 - 450 0.0480 0.0470 0.0420 0.0395 500 0.0480 0.0475 0.0420 0.0410

Rotation (rpm) 2.2% (w/) 2.2% (w/o) 3.5% (w/> 3.5% (w/o) 40 0.0060 - 0.0062 - 80 0.0090 0.0025 0.0073 0.0015 120 0.0140 - 0.0081 - 150 0.0170 0.0095 0.0099 0.0045 200 0.0215 - 0.0125 - 250 0.0250 0.0175 0.0138 0.0100 300 0.0300 - 0.0159 . 350 0.0320 0.0260 0.0170 0.0130 400 0.0325 - 0.0190 - 450 0.0325 0,0315 0.0220 0.0170 500 0.0320 0.0320 0.0230 0.0195 246

FIGURE 4.6(a): MAXIMAL kLa IN VARIOUS SYSTEMS STUDIED

Xanthan (g/L) STR-DT STR-MP STR-WIO CPBR-LC CPBR-GC 0.0 0.056 0.015 0.058 0.046 0.048 7.0 - 0.009 -- - 11.0 - 0.008 - -* 12.0 0.045 - 0.050 0.038 0.042 22.0 0.028 - 0.038 0.030 0.033 24.0 - 0.007 - - -

35.0 0.006 - 0.020 0.018 0.023

FIGURE 4.6(b): OTR IN VARIOUS SYSTEMS STUDIED

Xanthan (g/L) STR-DT STR-MP STR-WIO CPBR-LC CPBR-GC 0.0 1.26 0.34 1.30 1.03 1.09 7.0 - 0.18 ---

11.0 - 0.15 --- 12.0 0.86 - 0.95 0.72 0.78 22.0 0.49 - 0.66 0.52 0.62 24.0 - 0.12 - - - 35.0 0.09 - 0.34 0.31 0.38 APPENDIX E

EXPERIMENTAL DATA FOR CHAPTER V

FIGURE 5.2: ULTRAFILTRATION OF XANTHAN BROTH

(a) 0.84% Xanthan

Time (min) Flux Xanthan 0 5.80 0.84 1 7 5.56 - 28 5.45 - 4 1 5.56 - 54 5.60 - 71 5.40 - 87 5.00 - 1 1 1 4.75 - 132 5.00 - 158 4.92 - 185 5.50 - 209 5.08 1.39 231 5.45 - 261 5.44 - 291 6.14 - 319 6.23 2.15 342 7.08 - 361 8.60 - 383 7.40 3.74 397 7.75 4.47 425 5.78 6.38 439 5.78 - 454 4.20 10.41 462 3.33 11.78 470 3.00 12.60

247 248

(b) 3.37% Xanthan

Time (min) Flux Xanthan 0 6.20 3.37 1 0 6.00 20 6.40 24 6.23 3.75 52 5.80 80 6.18 5.00 108 5.78 136 5.79 156 5.10 8.63 1 66 4.20 1 76 3.60 9.96 184 3.20 194 3.10 1 1.20 216 2.90 13.13

(c) 4.47% Xanthan

Time (min) Flux Xanthan 0 6.00 4.47 20 6.10 - 37 6.29 5.30 74 5.78 6.38 1 1 2 5.78 8.02 138 5.00 9.54 152 4.30 10.38 165 3.65 1 1.23 173 3.20 1 1.73 189 3.00 12.80 249

FIGURE 5.4: EFFECT OF CELLS ON MEMBRANE FOUUNG

Cell-Free Heat-Treated Cell-Containing Time (min) Flux Flux Flux 0 1.85 1.85 1.82 1 0 1.86 1.83 1.83 20 1.84 1.85 1.82 30 1.87 1.86 1.81 40 1.86 1.83 1.76 50 1.87 1.82 1.72 60 1.86 1.82 1.64 70 1.86 1.81 1.59 80 1.84 1.82 1.57 90 1.82 1.80 1.56 100 1.86 1.81 1.54 1 1 0 1.85 1.82 1.54 1 20 1.82 1.80 1.53

FIGURE 5.6: EFFECT OF TRANSMEMBRANE PRESSURE

Water 2.5% Xanthan Trans Pressure Flux Rm Trans Pressure Flux Rf 0.69 25.50 2.70 0.69 1.64 3.93 1.17 46.04 2.55 1.24 1.82 6.57 1.72 71.76 2.40 1.79 2.10 8.29 250

FIGURE 5.7: XANTHAN CONCENTRATION ON VISCOSITY

500 f 1/S) 6 (1/s) Xanthan (%) Viscosity (cP) Xanthan (%) Viscosity fcP) 5.10 478.60 0.058 16.50 3.50 163.65 0.047 1 5.40 3.00 122.30 0.029 11.80 2.40 82.80 0.023 1 1.60 1.50 49.14 0.015 10.20 1.40 45.44 0.012 9.80 0.60 23.73 0.009 9.50 -- 0.006 9.20 - • 0.004 8.80

FIGURE 5.8: EFFECT OF XANTHAN CONCENTRATION

(a) Pump Efficiency

Xanthan % Pump 3.5 (in %) Pump 9.5 (in %) 0.0 100.00 100.00 0.8 98.21 - 2.7 94.09 97.78 5.2 87.54 95.28 8.4 81.24 92.48

10.8 77.12 -

(b) Filtrate Rate; (c) Jv and Rr

Xanthan % Filtrate Rate Jv Rf 10.03 4.20 1.07 15.87 6.55 6.50 1.66 10.13 2.72 7.10 1.82 9.22 0.84 15.00 3.84 4.24 0.20 37.50 9.61 1.55 0.13 62.50 16.02 0.84 0.06 96.80 24.81 0.45 0.03 205.00 52.54 0.09

0.00 390.00 - - 251 FIGURE 5.9: EFFECT OF SHEAR RATE ON XANTHAN VISCOSITY

Shear Rate O/sec) 3.6% Xanthan 1.8% Xanthan 0.9% Xanthan 0.01 1456160 182020 18202 0.02 1447840 180980 18098 0.04 1251680 156460 17646 0.06 996800 124600 15000 0.09 745360 93170 12460 0.15 532160 66520 11328 0.23 372672 46584 9 317 0.37 258776 32347 7258 0.59 177240 22155 475 8 0.94 101648 1 5206 3224 1.48 72624 10328 2179 2.35 47065 7024 1472 3.72 28790 4758 992 5.90 16774 3224 702 9.35 9716 2179 473 14.82 6887 1472 322 23.49 3967 992 218 37.23 2673 666 1 38 59.00 1801 450 95 93.52 1213 303 66 148.21 817 204 47 234.91 552 138 37 372.30 378 95 28 590.06 263 66 2 1 935.19 187 47 1 9

Shear Rate (1/s) 5.2% Xanthan 8.7% Xanthan 10.3% Xanthan 13.4% Xanthan 6.00 219 283 305 333 3.00 405 525 567 618 1.20 916 1 188 1283 1401 0.60 1698 2204 2381 2601 0.30 3147 4091 4418 4831 0.12 71 1 3 9264 10005 10949 0.06 13181 17191 18567 20332 0.03 24426 31902 34456 37758 252

FIGURE 5.10: EFFECT OF PUMPING RATE

Reading Q(ml/min) Shear Rate (1/s) Filtrate Rate Jv Rf 0.2 122.50 59.4 2.50 0.64 28.81 0.5 163.85 79.4 3.85 0.99 18.63 1.0 204.50 99.1 4.50 1.15 15.90 1.5 255.00 123.6 5.00 1.28 14.29 2.0 315.30 152.8 5.30 1.36 13.47 3.5 386.40 187.2 6.40 1.64 11.11 6.0 656.60 318.1 6.60 1.69 10.77 7.5 809.10 392.0 7.10 1.82 9.99 9.5 1011.60 490.1 7.60 1.95 9.32

FIGURE S.II: EFFECT OF pH

2.0% TIC 3.5% Lab Xan pH Viscosity (cP) pH Viscosity (cP) 3.95 1300 4.05 2252 5.09 1256 5.14 2288 6.02 808 6.34 2300 6.85 736 7.07 2288 8.40 960 8.10 2340 9.50 1 184 8.97 2408 10.60 1208 9.72 2428

Filtrate Rate (mL/min) Time (min) pH 4 pH 7 pH 10 0 5.80 5.80 6.00 1 0 5.80 5.80 5.80 20 5.70 5.90 5.80 30 5.80 5.80 5.80 40 5.70 5.80 5.70 50 5.60 5.70 5.70 60 5.70 5.70 5.60 70 5.70 5.70 5.70 80 5.80 5.70 5,70 90 5.70 5.60 5.60 100 5.70 5.70 5.70 110 5.80 5.70 5.70 120 5.80 5.70 5.70 253

FIGURE S. 12: RHEOLOGICAL PROPERTIES

rpm Before UF After UF 100.0 23.0 23.2 50.0 21.4 18.5 20.0 18.3 14.4 10.0 16.3 12.7 5.0 13.9 12.1 2.5 11.2 10.8 1.0 7.9 8.9 0.5 6.2 6.4

FIGURE 5.13: INTRINSIC VISCOSITY OF XANTHAN

Before UF After UF Xanthan (o/mL) U

EXPERIMENTAL DATA FOR CHAPTER VI

FIGURE 6.5: EFFECT OF XANTHAN CONCENTRATION ON K AND n

Xanthan (g/L) K n 0.0 -- 0.1 0.30 0.963 1.0 0.45 0.798 9.0 21 .59 0.288 18.0 121.84 0.200 27.0 355.22 0.142 36.0 776.18 0.1 13 52.0 1080.00 0.1 10 87.0 1398.00 0.108 103.0 1510.00 0.108 134.0 1649.00 0.107

FIGURE 6.6: EFFECT OF XANTHAN ON PRESSURE DROP

Pressure Drop (psig) Xanthan (g/L) Theoretical Experimental 0.1 - - 1.0 0.58 1.00 9.0 1.28 3.00 18.0 4.17 5.00 27.0 8.40 5.00 36.0 15.21 15.00 52.0 20.75 18.00 87.0 26.41 22.00 103.0 28.39 25.00 134.0 30.80 28.00

254 255

FIGURE 6.7: EFFECT OF PUMPING RATE ON PRESSURE DROP

Pressure Drop (psig) Pumping Rate (mL/s) Theoretical Experimental 10.12 19.66 18.00 8.09 19.19 17.50 6.57 18.75 1 8.00 3.86 17.69 17.00 3.15 17.30 16.00 2.55 16.90 15.50 2.05 16.49 14.00 1.64 16.09 14.00 1.23 15.59 12.00 1.00 15.24 - 0.60 14.41 - 0.30 13.35 - 0.10 1 1.83 - 0.05 10.96 - 0.03 10.37 - 0.01 9.19 -

0.00 0.00 -

FIGURE 6.8: EFFECT OF XA NTHA NON POWER CONSUMPTION

Xanthan (o/L) Power (W) F (mL/s) P/F fJ/mU 8.00 0.30 0.13 2.31 15.00 0.60 0.13 4.67 27.00 0.98 0.13 7.71 40.00 1.60 0.12 13.01 52.00 2.35 0.12 20.17 70.00 2.73 0.10 27.30 84.00 2.69 0.09 33.94 92.00 3.00 0.07 42.86 256 FIGURE 6.9: PUMPING RATE ON POWER CONSUMPTION

Q (mL/s) P (W) Filtrate Rate (mL/s) P/F (J/mL) 16.86 0.98 0.13 7.71 13.49 0.76 0.12 6.39 10.94 0.60 0.11 5.42 6.44 0.33 0.11 3.05 5.26 0.26 0.09 2.92 4.25 0.20 0.08 2.43 3.41 0.16 0.06 2.10 2.73 0.12 0.06 1.90 2.04 0.09 0.04 2.10

FIGURE 6.10: EFFECT OF APm ON POWER CONSUMPTION

Trans Pressure (Dsia) Power (W) Filtrate Rate (mL/s) P/F (J/mL) 5.00 0.54 0.07 7.45 10.00 0.54 0.09 5.99 25.00 0.54 0.12 4.55 40.00 0.54 0.13 3.98 55.00 0.54 0.15 3.64 70.00 0.54 0.16 3.40

FIGURE 6.11: OPERATING COSTS

$200/m 2 $400/m 2 Shear (1/s) P$/kg Mem$/kg Total ($/kg) Mem$/kg Total ($/kg) 490.15 1.42 7.34 8.76 1.47 2.89 392.03 0.98 7.88 8.86 1.56 2.56 318.14 0.83 8.43 9.26 1.69 2.52 187.22 0.55 9.98 10.53 2.00 2.54 152.77 0.47 10.66 11.12 2.13 2.60 123.55 0.39 1 1.40 1 1.80 2.28 2.68 99.09 0.33 12.24 12.57 2.45 2.78 79.39 0.28 13.14 13.42 2.63 2.91 59.35 0.22 14.42 14.64 2.88 3.10 FIGURE 6.12: OPERATING COSTS

Select Shear 153 (1/s) P (10M Pa) Mem $/kg Energy $/kg Total $/kg 3.45 2.06 0.90 2.96 6.89 1 .73 0.76 2.49 17.24 1 .38 0.60 1.98 27.58 1.23 0.54 1.76 37.92 1.13 0.50 1.63 48.26 1.07 0.47 1 .53

FIGURE 6.13: ENERGY CONSUMPTION

Final Xanthan UF (40L), J Alcohol (KJ) Total (KJ) 40.00 79581.21 5 7 8 2 4 5 .0 0 578324.58 55.00 2 4 5 7 3 2 .4 5 420 5 4 1 .8 2 4207 8 7 .5 5 70.00 443453.24 330425.71 3 30869.17 85.00 660 6 0 0 .3 4 272 1 1 5 .2 9 272 7 7 5 .8 9 1 00.00 887 0 1 2 .8 4 231 2 9 8 .0 0 232185.01 1 15.00 1122278.84 2011 2 8 .7 0 202250.97 130.00 1360439.40 177921.54 179281.98 150.00 1667254.78 154198.67 155865.92 25.00 0.00 925 1 9 2 .0 0 925192.00