Membrane Technology: Volume 3: Membranes for Food Applications

Edited by Klaus-Viktor Peinemann, Suzana Pereira Nunes, and Lidietta Giorno Membrane Technology Related Titles Peinem...

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Edited by Klaus-Viktor Peinemann, Suzana Pereira Nunes, and Lidietta Giorno Membrane Technology

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Membrane Technology Volume 3: Membranes for Food Applications

Edited by Klaus-Viktor Peinemann, Suzana Pereira Nunes, and Lidietta Giorno

The Editors Dr. Klaus-Viktor Peinemann KAUST Membranes Research Center Bldg.2, 3rd Floor,Room 3216-W9 4700 King Abdullah Univ. 23955-6900 Thuwal Saudi Arabia Dr. Suzana Pereira Nunes KAUST Membranes Research Center Bldg.2, 3rd Floor,Room 3216-W9 4700 King Abdullah Univ. 23955-6900 Thuwal Saudi Arabia Prof. Lidietta Giorno University of Calabria Institute on Membrane Technology Via P. Bucci 17/C 87036 Rende (CS) Italia

All books published by Wiley-VCH are carefully produced. Nevertheless, authors, editors, and publisher do not warrant the information contained in these books, including this book, to be free of errors. Readers are advised to keep in mind that statements, data, illustrations, procedural details or other items may inadvertently be inaccurate. Library of Congress Card No.: applied for British Library Cataloguing-in-Publication Data A catalogue record for this book is available from the British Library. Bibliographic information published by the Deutsche Nationalbibliothek The Deutsche Nationalbibliothek lists this publication in the Deutsche Nationalbibliograﬁe; detailed bibliographic data are available on the Internet at http://dnb.d-nb.de. # 2010 WILEY-VCH Verlag GmbH & Co. KGaA, Boschstr. 12, 69469 Weinheim All rights reserved (including those of translation into other languages). No part of this book may be reproduced in any form – by photoprinting, microﬁlm, or any other means – nor transmitted or translated into a machine language without written permission from the publishers. Registered names, trademarks, etc. used in this book, even when not speciﬁcally marked as such, are not to be considered unprotected by law. Cover Design Adam Design, Weinheim Typesetting Thomson Digital, Noida, India Printing and Binding betz-druck GmbH, Darmstadt Printed in the Federal Republic of Germany Printed on acid-free paper ISBN: 978-3-527-31482-9

V

Contents Preface XI List of Contributors

1 1.1 1.2 1.2.1 1.2.2 1.2.2.1 1.2.2.2 1.2.2.3 1.2.2.4 1.2.2.5 1.3 1.3.1 1.3.1.1 1.3.1.2 1.3.1.3 1.3.2 1.3.2.1 1.3.2.2 1.3.2.3 1.3.2.4 1.3.3 1.3.3.1 1.4 1.4.1 1.4.2 1.5

XIII

Cross-Flow Membrane Applications in the Food Industry 1 Frank Lipnizki Introduction 1 Dairy Industry 2 Dairy Industry Overview 2 Key Membrane Applications 3 Removal of Bacteria and Spores from Milk, Whey and Cheese Brine 3 Milk Protein Standardization, Concentration and Fractionation 5 Whey Protein Concentration and Fractionation 5 Whey Demineralization 7 Cheese Manufacturing 8 Fermented Food Products 9 Beer 9 Beer from Tank Bottoms/Recovery of Surplus Yeast 9 Beer Clariﬁcation 11 Beer Dealcoholization 11 Wine 12 Must Correction 13 Clariﬁcation of Wine 13 Rejuvenation of Old Wine (Lifting) 13 Alcohol Removal 14 Vinegar 14 Clariﬁcation of Vinegar 15 Fruit Juices 15 Fruit-Juice Clariﬁcation 16 Fruit-Juice Concentration 17 Other Membrane Applications in the Food Industry 17

VI

Contents

1.5.1 1.5.2 1.6 1.6.1 1.6.2 1.6.2.1 1.6.2.2 1.6.2.3 1.6.3

Membrane Processes as Production Step 18 Membrane Processes for Water and Wastewater 18 Future Trends 18 New Applications of Membrane Processes 20 New Membrane Processes 20 Pervaporation 21 Electrodialysis 21 Membrane Contactors – Osmotic Distillation 22 Integrated Process Solutions: Synergies and Hybrid Processes References 23

2

Membrane Processes for Dairy Fractionation 25 Karin Schroën, Anna M.C. van Dinther, Solomon Bogale, Martijntje Vollebregt, Gerben Brans, and Remko M. Boom Introduction 25 Membrane Separation of Components 27 Removal of Milk Fat from Whole Milk 27 Removal of Bacteria and Spores from Skim Milk (Cold Pasteurization) 27 Concentration of Casein Micelles in Skim Milk 28 Recovery of Serum Proteins from Cheese Whey 30 Methods to Enhance Membrane Separation 30 Critical Flux Concept 31 Uniform Low Transmembrane Pressure Concept (UTP) 32 Turbulence Promotion 32 Backpulsing and Flow Reversal 33 Other Methods 33 Use of Models for Membrane Separation 34 How to Get from Separation to Fractionation 35 Membranes with Uniform Pore Size 36 Simulation of Particle Behavior 37 Membrane Modiﬁcation 37 Outlook 38 References 39

2.1 2.2 2.2.1 2.2.2 2.2.3 2.2.4 2.3 2.3.1 2.3.2 2.3.3 2.3.4 2.3.5 2.4 2.5 2.5.1 2.5.2 2.5.3 2.6

3

3.1 3.1.1 3.1.1.1 3.1.1.2 3.2 3.2.1 3.2.1.1

23

Milk and Dairy Effluents Processing: Comparison of Cross-Flow and Dynamic Filtrations 45 Michel Y. Jaffrin, Valentina S. Espina, and Matthieu Frappart Introduction 45 Properties and Applications of Various Proteins 45 Caseins 45 Whey Proteins 45 Applications of Membrane Cross-Flow Filtration to Milk Processing Milk Microﬁltration 46 Bacteria and Spore Removal 46

46

Contents

3.2.1.2 3.2.2 3.2.2.1 3.2.2.2 3.2.3 3.2.3.1 3.2.3.2 3.3 3.3.1 3.3.2 3.3.3 3.3.3.1 3.3.3.2 3.3.3.3 3.3.3.4 3.3.4 3.4

Casein Micelles Separation from Whey Proteins 47 Milk Ultraﬁltration (UF) 48 Total Proteins Concentration 48 Whey-Protein Fractionation 49 Applications of Milk Nanoﬁltration (NF) and Reverse Osmosis (RO) 50 Treatment of Cheese Whey and Fabrication of Yogurts 50 Treatment of Dairy Efﬂuents 51 Dynamic Filtration 51 Principle and Advantages of Dynamic (Shear-Enhanced) Filtration 51 Industrial Dynamic Filtration Systems 52 Application of Dynamic Filtration to Skim-Milk Processing 55 Casein Separation from Whey Proteins by MF 55 Dynamic Ultraﬁltration of Skim Milk 60 Total Protein Concentration by UF for Cheese Manufacturing 63 a-La and b-Lg Protein Fractionation by UF 64 Treatment of Dairy-Process Waters by Dynamic NF and RO 66 Conclusion 69 References 70

4

Electrodialysis in the Food Industry 75 Jamie Hestekin, Thang Ho, and Thomas Potts Introduction 75 Technology Overview 76 Principle of the Electrodialysis Process 76 System Design 79 Concentration Polarization, Limiting Current Density, Current Utilization, and Power Consumption 79 System Design and Cost Analysis 80 Electrodialysis Applications in the Food Industry 82 Wine Industry 83 Juice and Sugar Industry 85 Dairy Industry 88 Protein Fractions 94 Hybrid Technologies 96 Electrodeionization 96 Electrochemical Coagulation 97 Electroreduction 98 Conclusion and Future Innovations 98 References 101

4.1 4.2 4.2.1 4.2.2 4.2.2.1 4.2.2.2 4.3 4.3.1 4.3.2 4.3.3 4.3.4 4.4 4.4.1 4.4.2 4.4.3 4.5

5 5.1 5.2 5.3

Membrane Processes in Must and Wine Industries 105 Maria Norberta De Pinho Introduction 105 Wine Clariﬁcation by Microﬁltration and Ultraﬁltration 106 Wine Tartaric Stabilization by Electrodialysis 111

VII

VIII

Contents

5.4 5.5

6 6.1 6.1.1 6.1.2 6.2 6.2.1 6.2.2 6.3 6.4

7

7.1 7.1.1 7.1.2 7.1.3 7.1.4 7.1.5 7.2 7.2.1 7.2.2 7.2.3 7.3 7.3.1 7.3.1.1 7.3.1.2 7.3.2 7.3.2.1 7.3.2.2 7.3.3 7.3.3.1 7.3.3.2 7.4 7.4.1

Inﬂuence of MF/UF Polysaccharide Removal on Wine Tartaric Stability 113 Nanoﬁltration of Grape Must for Sugar/Organic Acids Fractionation 115 References 117 New Applications for Membrane Technologies in Enology 119 Martine Mietton Peuchot Reduction of Alcohol Content 119 Reduction of Must Sugars to Obtain a Lower Alcohol Content in Wines 119 Reduction of Alcohol Content in Wine 121 Reduction of Malic Acid in Grape Musts or Volatile Acidity in Wines 123 Reduction of Malic Acid in Musts 123 Reduction of Volatile Acidity 123 Acidiﬁcation of Musts and Wines 125 Other Potential Applications 126 References 127 Membrane Emulsification for Food Applications 129 Henelyta S. Ribeiro, Jo J. M. Janssen, Isao Kobayashi, and Mitsutoshi Nakajima Introduction 129 Cross-Flow Membrane Emulsiﬁcation (XME) 130 Dead-End Membrane Emulsiﬁcation (PME) 131 Rotating Membrane Emulsiﬁcation (RME) 131 Vibrating Membrane Emulsiﬁcation (VME) 132 Microchannel Emulsiﬁcation 133 Understanding of the Process at the Pore Level 134 XME 134 PME 137 MCE 138 Production of Structured Systems for Food Applications 140 O/W Emulsions 140 Membrane Emulsiﬁcation 140 Microchannel Emulsiﬁcation 141 W/O Emulsions 144 Membrane Emulsiﬁcation 144 Microchannel Emulsiﬁcation 145 W/O/W Emulsions 147 Membrane Emulsiﬁcation 147 Microchannel Emulsiﬁcation 149 Encapsulation of Active Molecules 149 Membrane Emulsiﬁcation 149

Contents

7.4.2 7.5 7.5.1 7.5.1.1 7.5.1.2 7.5.1.3 7.5.2 7.5.2.1 7.5.2.2 7.5.2.3 7.5.2.4 7.5.3 7.5.3.1 7.5.3.2 7.5.4 7.6

8

8.1 8.2 8.3 8.3.1 8.3.2 8.3.3 8.3.4 8.4 8.4.1 8.4.2 8.4.3 8.4.4 8.4.5 8.5 8.5.1 8.5.2 8.6

9 9.1

Microchannel Emulsiﬁcation 151 Assessment of the Potential Beneﬁts of Membrane Emulsiﬁcation in Foods 151 DSD and Product Stability 152 Physical Stability 152 Chemical Stability 153 Microbiological 153 DSD and Product Rheology 153 Yield Stress 154 Elastic Modulus 155 Viscosity 155 Formation of Flocculated Networks 155 Product Properties Related to Low-Shear Processing 156 Shear Damage to Ingredients 156 Effect on Product Structure 157 Summary 157 Conclusions 158 References 158 Membrane Contactors in Integrated Processes for Fruit-Juice Processing 167 Alfredo Cassano and Enrico Drioli Introduction 167 Membrane Contactors: Fundamentals 169 Osmotic Distillation 171 Process Fundamentals 171 OD Membranes and Modules 172 Effect of Operating Conditions on the OD Flux 174 OD Applications 177 Membrane Distillation 183 Process Fundamentals 183 MD Membranes and Modules 184 Effect of Operating Parameters on MD Fluxes 186 MD Applications 187 Coupled Operation of Osmotic Distillation and Membrane Distillation 190 Supported-Liquid Membranes 191 Process Fundamentals 191 SLMs Applications 192 Conclusions 193 References 195 Membrane Bioreactors in Functional Food Ingredients Production 201 Rosalinda Mazzei, Sudip Chakraborty, Enrico Drioli, and Lidietta Giorno Introduction 201

IX

X

Contents

9.2 9.3 9.4 9.5 9.5.1 9.5.2 9.5.3 9.6 9.6.1 9.6.1.1 9.6.1.1.1 9.6.1.1.2 9.7

10 10.1 10.2 10.2.1 10.3 10.4

Membrane Bioreactors and Functional Food 201 Membrane Bioreactor in Sugar and Starch Processing 203 Membrane Bioreactor in Oil and Fat Processing Industry 206 Membrane Bioreactors in Hard Drink Industry and Liquid Beverages 208 Wine 208 Beer 212 Ethanol Production 212 Membrane Bioreactor in Other Liquid Beverages 213 Fruit-Juices Production 213 Functional Food Production in the Milk and Whey Field by Membrane Bioreactor 214 Lactose Hydrolysis 215 Protein Hydrolysis in Milk and Whey by MBR 215 Conclusions 219 References 219 Membranes for Food Packaging 223 Alberto Figoli, Erika Mascheroni, Sara Limbo, and Enrico Drioli Introduction 223 Application of Membranes in Controlling Gas Permeability 225 Membranes in Modiﬁed-Atmosphere Packaging 228 Membranes as Devices for Active Food Packaging 232 Conclusions 238 References 239 Index

241

XI

Preface In the Membrane Technology book series we collect and present in different volumes the most relevant examples of how synthetic membranes are contributing to ﬁnding solutions to key issues of the world population. We cover essential topics starting with life science, followed by energy and water. In this volume, we also approach certainly one of the most crucial aspects for everyday life: food. Membranes have been shown very early to be useful in several processes of food industries. The dairy industry was one of the ﬁrst sectors to proﬁt from membrane technology on a large scale. Nowadays, a large part of the available milk products in developed countries involves at least one step using membrane processes. Whey processing and cheese manufacture are good examples. Membranes can make the processes more effective and bring quality advantages, which are hardly beaten by traditional methods. In recent decades membranes have become a routine technology also in other food industries. The needs for transportation at long distances have stimulated the use of membranes to concentrate juices. Membranes have been the technology of choice in applications where keeping aroma and processing at mild temperatures is essential. It has led to new process routes and to reducing droplet sizes in emulsiﬁcation techniques. The market for nonalcoholic beer is growing and it is still a big challenge to keep the taste like the original products. Membranes are substituting steps of manufacture of the most traditional industries, like wine production. Finally, membranes play an essential role also in food packaging, where concepts of gas permeability are important to meet the new demands of food safety and storage. This volume will appeal to workers in the ﬁeld of membrane technology applied to food, bringing together information on the already established and the potential technologies in this ﬁeld. Thuwal, Kingdom of Saudi Arabia

Dr. Klaus-Viktor Peinemann

XIII

List of Contributors Solomon Bogale Kassa Addis Ababa University, Faculty of Technology Department of Chemical Engineering P.O. Box 385 Addis Ababa Ethiopia and Wageningen University, Dept. Agrotechnology and Food Sciences Bomenweg 2 Postbus 8129 6703 HD, Wageningen The Netherlands Remko M. Boom Wageningen University, Dept. Agrotechnology and Food Sciences Bomenweg 2 Postbus 8129 6703 HD, Wageningen The Netherlands Gerben Brans Applikon De Brauwweg 13 3125 AE Schiedam The Netherlands

Alfredo Cassano Institute on Membrane Technology National Research Council ITM-CNR Via P. Bucci 17/C 87036 Rende (CS) Italy Anna M. C. van Dinther Wageningen University Dept. Agrotechnology and Food Sciences Bomenweg 2 Postbus 8129 6703 HD, Wageningen The Netherlands Enrico Drioli Institute on Membrane Technology National Research Council ITM-CNR Via P. Bucci 17/C 87036 Rende (CS) Italy Valentina S Espina Olygose SAS Parc Technologique de lOise BP 50149 60201 Venette France

XIV

List of Contributors

Alberto Figoli Institute on Membrane Technology National Research Council ITM-CNR Via P. Bucci 17/C 87036 Rende (CS) Italy

Isao Kobayashi National Food Research Institute, NARO Food Engineering Division 2-1-12 Kannondai 305-8642 Tsukuba, Ibaraki Japan

Matthieu Frappart GEPEA-CRTT, Faculte des Sciences et Techniques BP 406 44602 Saint Nazaire Cedex France

Sara Limbo Department of Food Science and Microbiology (DISTAM) University of Milan Via Celoria 2 20133 Milan Italy

Lidietta Giorno Institute on Membrane Technology National Research Council ITM-CNR Via P. Bucci 17/C 87036 Rende (CS) Italy

Frank Lipnizki Business Centre Membranes Alfa Laval Copenhagen A/S Maskinvej 5 DK-2860 Sborg, Denmark

Jamie Hestekin Ralph E. Martin Department of Chemical Engineering 3202 Bell University of Arkansas Fayetteville, AR 72701-1201 USA

EriKa Mascheroni Department of Food Science and Microbiology (DISTAM) University of Milan Via Celoria 2 20133 Milan Italy

Michel Y. Jaffrin Biological Engineering Technological University of Compiegne UMR 6600 EA 4297 TIMR, BP 20529 60205 Compiegne France

Mitsutoshi Nakajima University of Tsukuba Graduate School of Life and Environmental Sciences 1-1-1 Tennodai, 305-8572 Tsukuba, Ibaraki Japan

Jo J. M. Janssen Unilever R&D Vlaardingen Olivier van Noortlaan 120 3133 AT Vlaardingen The Netherlands

List of Contributors

Martine Mietton Peuchot Université Bordeaux ISVV, UMR 1219 Oenology Faculty of Enology 210 Chemin de Leysotte CS50008, 33882, Villenave dOrnon France Maria Norberta De Pinho Department of Chemical and Biological Engineering Instituto Superior Técnico Technical University of Lisbon Av. Rovisco Pais-1 1049-001 Lisboa Portugal Henelyta Santos Ribeiro Unilever R&D Vlaardingen Olivier van Noortlaan 120 3133 AT Vlaardingen The Netherlands

Karin Schroën Wageningen University Dept. Agrotechnology and Food Sciences Bomenweg 2 Postbus 8129 6703 HD, Wageningen The Netherlands Martijntje Vollebregt Wageningen University and Research Centre Department of Agrotechnology and Food Sciences A&F institute, Fresh, Food and Chains Bornse Weilanden 9 6708 WG Wageningen The Netherlands

XV

j1

1 Cross-Flow Membrane Applications in the Food Industry Frank Lipnizki 1.1 Introduction

Over the last two decades, the worldwide market for membrane technology in the food industry increased to a market volume of about D 800–850 million and is now the second biggest industrial market for membranes after water and wastewater treatment including desalination. The key membrane technologies in the food industry are the pressure-driven membrane processes microﬁltration (MF), ultraﬁltration (UF), nanoﬁltration (NF) and reverse osmosis (RO). The market share of UF systems and membranes accounts for the largest share of the membrane market with 35%, followed by MF systems and membranes with a share of 33%, and NF/RO systems and membranes with a share of 30%. Other membrane processes such as membrane contactors (MC), electrodialysis (ED) and pervaporation (PV) have only a small market share. The major applications in this market are in the dairy industry (milk, whey, brine, etc.) followed by other beverage industries (beer, fruit juices, and wine, etc.). The success of membrane technology in the food and beverage market is directly linked to some of the key advantages of membrane processes over conventional separation technologies. Among these advantages are: . . . .

gentle product treatment due to moderate temperature changes during processing; high selectivity based on unique separation mechanisms, for example sieving, solution-diffusion or ion-exchange mechanism; compact and modular design for ease of installation and extension; low energy consumption compared to condensers and evaporators.

The key disadvantage of membrane ﬁltration is the fouling of the membrane causing a reduction in ﬂux and thus a loss in process productivity over time. The effect of fouling can be minimized by regular cleaning intervals. In the food industry it is common to have at least one cleaning cycle per 24-h shift. Other actions to reduce fouling are directly related to plant design and operation. During the plant design, the selection of a low-fouling membrane, for example hydrophilic membranes to reduce fouling by bacteria, and membrane modules with appropriate channel heights,

j 1 Cross-Flow Membrane Applications in the Food Industry

2

for example modules with open channel design to avoid blockage by particles, can reduce the risk of fouling and contamination signiﬁcantly. Operating the plant below the critical ﬂux – the ﬂux below which a decline of ﬂux over time does not occur, and above which fouling is observed – can extend the time between cleaning intervals signiﬁcantly but is commonly related to low-pressure/low-ﬂux operation, which translates into low capacities. Alternatively, operating the process in turbulent ﬂow regime can reduce the effect of fouling, but the generation of turbulence is linked to an increase in pressure drop and therefore higher energy costs. Other limitations to the application of membrane processes might be related to the feed characteristics, for example increase of viscosity with concentration, or to separation mechanisms used in the membrane process, for example osmotic pressure increases with concentration. In the following, successful applications of membrane processes in the food industry will be introduced. The ﬁrst part of this chapter will focus on the dairy industry, the largest and most developed membrane market in the food industry, followed by the fermented food products – beer, wine and vinegar – fruit juices and other established membrane applications. The ﬁnal section of this chapter will give an outlook of potential membrane applications in the food industry focusing especially on the emerging membrane technologies: membrane contactors, pervaporation and electrodialysis.

1.2 Dairy Industry 1.2.1 Dairy Industry Overview

The dairy industry has used membrane processing since its introduction in the food industry in the late 1960s to clarify, concentrate and fractionate a variety of dairy products. Applying membrane technology to whey processing allowed the production of reﬁned proteins and commercial usage and thus transformed a waste byproduct from cheese production into a valuable product. In addition to whey processing, membrane technology is also used for ﬂuid milk processing with clear advantages. Further, speciﬁc milk components can be obtained without causing a phase change to the ﬂuid milk by the addition of heat as in evaporation, or an enzyme, as done in most cheese-making techniques. The ﬁltered milk can then be directly used in the manufacture of such dairy products as cheese, ice cream and yoghurt. By applying membranes with different pore sizes and molecular weight cut-offs (MWCOs), the milk can be modiﬁed by separating, clarifying, or fractionating a selected component in milk from other components. The pressure-driven membrane processes MF, UF, NF and RO are the most common membrane processes in the dairy industry and based on their applicability range it is possible to separate virtually every major component of milk as shown in Figure 1.1, thus enabling the manufacturing of products with unique properties and functionalities.

1.2 Dairy Industry

Figure 1.1 Milk processing with membrane technology.

1.2.2 Key Membrane Applications

In the following, the key applications of cross-ﬂow membrane technology in the dairy industry are discussed. 1.2.2.1 Removal of Bacteria and Spores from Milk, Whey and Cheese Brine The removal of bacteria and spores from milk to extend its shelf-life by MF is an alternative way to ultrapasteurization. In this approach, the organoleptic and chemical properties of the milk are unaltered. The ﬁrst commercial system of this so-called Bactocatch was developed by Alfa Laval [1–3] and marketed by Tetra Pak under the name Tetra AlcrossÒ Bactocatch. In this process, the raw milk is separated into skim milk and cream, see Figure 1.2. The resulting skim milk is microﬁltered using ceramic membranes with a pore size of 1.4 mm at constant transmembrane pressure (TMP). Thus, the retentate contains nearly all the bacteria and spores, while the bacterial concentration in the permeate is less than 0.5% of the original value in milk. The retentate is then mixed with a standardized quantity of cream. Subsequently, this mix is subjected to a conventional high heat treatment at 130 C for 4 s and reintroduced into the permeate, and the mixture is then pasteurized. Since less than 10% of the milk is heat treated at the high temperature, the sensory quality of the milk is signiﬁcantly improved.

j3

j 1 Cross-Flow Membrane Applications in the Food Industry

4

Figure 1.2 Bacterial removal from milk by MF.

MF for the removal of bacteria and spores can be further applied in the production of other dairy products. In the production of cheese, the use of low bacterial milk improves also the keeping quality of cheese due to the removal of spores, thus eliminating the need of additives (e.g., nitrate). While in the production of whey protein concentrates (WPC) and isolates (WPI), this MF concept is used to remove bacteria and spores giving a high quality product (see Figure 1.4). Hence, by applying MF the heat treatment of the WPC/WPI is kept to a minimum, which preserves the functional properties of the whey proteins. Finally, in the manufacture of cheese the concentrated curd is submerged in a salt solution to improve the cheese preservation and to develop the ﬂavor and other cheese properties. This process is called brining. Efﬁcient sanitation of cheese brine has become a major concern to the dairy industry in recent years. This results from the possibility of post-contamination of cheeses in the brine, especially by pathogenic bacteria. The application of MF for sanitation of cheese brine, using ceramic or spiralwound membranes, results in a superior cheese quality compared to the traditional processes of heat treatment and kieselguhr ﬁltration. MF has the advantages of being simple to perform, of maintaining the chemical balance of the brine and of eliminating ﬁlter aids. In the brine treatment by MF it is normally necessary to make a preﬁltration of the brine solution, which is easily done by dead-end ﬁlter bag or cartridge with a pore size of 100 mm [4].

1.2 Dairy Industry

1.2.2.2 Milk Protein Standardization, Concentration and Fractionation The protein content of milk is subjected to natural variations during the year. Standardization of milk by UF offers the possibility of increasing or decreasing the protein content in milk without the need of adding milk powders, casein and whey protein concentrates. Skim milk and 1% milk with increased protein content have an improved appearance (whiter milk) and higher viscosity [5]. The sensory quality of increased protein milk is therefore more similar to that of higher fat milks resulting in an improved consumer appeal. Another application of UF is the standardization of protein and total solids in milk for use in fermented dairy products, such as cream cheeses, yoghurt and cottage cheeses. The resulting dairy products have superior quality and sensory characteristics compared to those produced from milk concentrated by conventional methods [6]. With the quality obtained by membrane ﬁltration, attributes such as consistency, post-processing and extent of syneresis are easier to control. However, the use of membrane-processed milk often requires an adjustment in starter culture selection and fermentation conditions due to the compositional changes in the UF milk. Concentration of milk, which conventionally is done by evaporation techniques, can also be achieved by RO. The concentrated milk has its greatest potential in ice-cream manufacturing, since all the solids are retained in the concentrate and 70% of the water is removed. MF and/or UF are used in the production of milk protein concentrates (MPC), which are products containing 50–58% of protein. These products are used as food additives and it is therefore extremely important to maintain the functionality of the proteins. By using UF membranes in combination with MF and/or diaﬁltration (DF) with the corrected adjustments of pH, temperature and ﬁltration conditions, it is possible to produce the desirable MPC for a speciﬁc food application. The most promising MFapplication in the dairy industry is the fractionation of milk protein. The separation of micellar casein from the whey proteins can be achieved by ceramic membranes with a pore size of 0.2 mm at a constant TMP. The resulting retentate has a high concentration of native calcium phosphocaseinate that can be used for cheese making. Native casein has an excellent rennet-coagulation ability that will make calcium phosphocaseinate an exceptional enrichment for cheese-milk. The permeate can be further processed by UF to produce high-quality WPC. These protein concentrates can be further separated into lactoferrin, b-lactoglobulin and a-lactalbumin via ion-exchange chromatography. Both b-lactoglobulin and a-lactalbumin have great potential markets. b-lactoglobulin can be used as a gelling agent and a-lactalbumin, which is very rich in tryptophan, can be used in the production of peptides with physiological properties. Another application can be the production of infant milk. The fractionation of milk proteins using membrane technology enables the recovery of value-added protein ingredients. Further, the casein and whey proteins are separated without the need of heat or enzymes. The potential applications of membrane separation in milk processing are shown in Figure 1.3. 1.2.2.3 Whey Protein Concentration and Fractionation Whey is a by-product from the cheese industry. It has low content of solids and high biological oxygen demand (BOD), which creates a major disposal problem for the

j5

j 1 Cross-Flow Membrane Applications in the Food Industry

6

Figure 1.3 Applications of membrane technology in milk processing.

dairy industry. In the past, all whey was disposed of as sewage, sprayed on ﬁelds or used for animal feed. By applying membrane technology whey can be concentrated to produce WPC and WPI, as well as fractionated and puriﬁed to obtain puriﬁed a-lactalbumin and b-lactoglobulin. Hence, a once wasted product can be converted into high value-added products and at the same time one of the key pollution problems of the dairy industry can be solved. Consequently, the use of UF and RO to concentrate whey was one of the ﬁrst applications of membranes in the dairy industry. Due to the complexity and diversity of whey, it is necessary to use different membrane processes to produce a speciﬁc product (see Figure 1.4). The production of WPC with 35–85% protein in the total solids can be achieved by a combination of UF and DF. MF can be used as a pretreatment to remove both bacteria and fat and allows the production of WPI with 90% protein in the total solids. Whey proteins have not only a high nutritional value but also functional properties. They can be used as gelling, emulsifying and foaming agents. Therefore, whey concentrates have farreaching applications not only in dairy foods, but also in confectionary, nutritional foods, beverages and even processed meats. The presence of fat in whey leads to decreased functional properties and shorter storage time. Several processes involving membranes have been developed to remove the residual fat from whey [7–11]. The most common process, developed by Maubois et al. [9] and Fauquant et al. [8], exploits the ability of the phospholipids to aggregate by calcium binding under moderate heat treatment for 8 min at 50 C. This process is called thermocalcic precipitation. Defatted whey is then obtained by MF with a pore size of 0.14 mm to separate the resulting precipitate. Defatted whey can be further processed by UF, which also improves the performance in the subsequent membrane processes. The defatted WPC has a foaming capacity similar to that of egg white and the same protein content. Its applications can be as raw material in the pastry and

1.2 Dairy Industry

Figure 1.4 Applications of membrane technology in whey processing.

icecream production. The MF retentate, which contains a high amount of phospholipids, can be used as an effective emulsiﬁer agent for food and cosmetic applications. The puriﬁed proteins b-lactoglobulin and a-lactalbumin can be obtained from the defatted whey. At low pH (4.0–4.5) and under moderate heat treatment for 30 min at 55 C, a-lactalbumin polymerizes reversibly entrapping most of the residual lipids and the other whey proteins with the exception of the b-lactoglobulin. The fractionation of b-lactoglobulin from the remaining proteins can then be done by MFwith a pore size of 0.2 mm or centrifugation. The resulting soluble phase, rich in b-lactoglobulin, can be further puriﬁed by UF coupled with electrodialysis (ED) or DF [9]. Puriﬁcation of a-lactalbumin from the MF retentate can be achieved by solubilization at a neutral pH and subsequently by UF using a membrane with an MWCO of 50 000 Dalton. It has also been reported that membranes can be applied for the isolation of Kcasein-glycomacropeptid (GMP) from cheese whey. GMP can ﬁnd several applications in the pharmaceutical industry. Studies have shown that GMP avoids the adhesion of Escherichia coli cells to the intestine walls, protects against inﬂuenza and prevents adhesion of tartar to teeth [12]. It should also be noted that membrane ﬁltration also plays a major role in the lactose manufacture from whey using UF and RO and in the production of lowcarbohydrate beverages with high dairy protein content. 1.2.2.4 Whey Demineralization In the dairy industry, the NF process is used to concentrate and partially demineralize liquid whey. Due to the selectivity of the membranes most of the monovalent ions, the

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organic acids, and some of the lactose will pass the membrane. NF is a very interesting alternative to ion exchange and ED if moderate demineralization is required. One advantage of NF compared to the other two processes is that NF is a simple process, which partially demineralizes and concentrates the whey at the same time. The maximum level of demineralization by NF is about 35% reduction of the ash content with a concentration factor of about 3.5–4. By applying a DF step it is possible to increase the level of demineralization up to 45%. Other applications of NF in whey processing include: concentration and partial demineralization of whey UF permeates prior to the manufacture of lactose and lactose derivatives, converting salt whey to normal whey while solving a disposal problem, treating cheese brine solutions to be reused. The potential applications of membrane separation in whey processing are shown in Figure 1.4. 1.2.2.5 Cheese Manufacturing Another early application of membrane technology in the dairy industry was in cheese manufacturing for production of Feta cheese and brine treatment by UF. Nowadays, membrane-processed milk is also successfully used in the manufacturing of quark and cream cheeses. Together with WPC production, the use of UF milk for the production of cheese is the most widespread application of membranes in the dairy industry. The advantages of UF concentrated milk in cheese making compared to traditional methods are the following: . . . . .

increases the total solids, which increases the cheese yield and therefore decreases the production costs in terms of energy and equipment; reduces the rennet and starter culture requirements since UF-milk has a good ability of enzymatic coagulation; reduces the wastewater processing costs of the cheese plant; improves the quality and composition control; increases the nutritional value due to the incorporation of the whey protein in the cheese. UF in cheese processing can be used in three ways [6]:

Preconcentration – The standardized cheese milk is concentrated by a factor of 1.2–2 and it can be used for most cheese types. This allows the capacity of the cheese vats and whey draining equipment to be doubled. However, the cheese yield will not be signiﬁcantly improved since only 4.5–5% of the protein content is increased. It is used to produce Cheddar, Cottage Cheese and Mozzarella, and it can be used to standardize cheese milk and manipulate its mineral composition, resulting in a more consistent quality in the ﬁnal product. 2) Partial concentration – The standardized cheese milk is concentrated by a factor 2–6. It is used in the manufacture of Cheddar cheese by using for example, the APV-SiroCurd process, in which the milk is concentrated ﬁve times with DF in order to standardize the salt balance [13]. It is also used to produce other cheese types like Queso Fresco, structure Feta, Camembert and Brie. 1)

1.3 Fermented Food Products

3)

Total concentration – The standardized cheese milk is concentrated to the total solids content in the ﬁnal cheese. This provides the maximum yield increase and since there is no whey drainage, the cheese can be manufactured without the need for a cheese vat. It is used to produce cast Feta, quark, cream cheese, Ricotta and Mascarpone.

The UF permeate, which contains mainly lactose, can be concentrated by RO. The permeate from the RO process can be polished by another RO unit. After pasteurization or UV light treatment, the permeate from the polisher can be used at the plant as process water, thus reducing the water costs of the plant. Although UF has advantages in cheese production, the increase of whey content in the cheese due to the concentration of all milk proteins can have a negative effect on the ripening of semihard and hard cheeses [14, 15]. Therefore, UF should be viewed as a complementary process to cheese manufacturing and not as an alternative process.

1.3 Fermented Food Products

In the production of the fermented food products, for example beer, wine and vinegar, membranes have initially established themselves as a clariﬁcation step after the fermentation. Initially, dead-end ﬁlters were used in the production of fermented food products followed by the ﬁrst trials of cross-ﬂow ﬁltration for the clariﬁcation of beer, wine and vinegar in the 1970s. However, the ﬁrst industrial application in this segment was the dealcoholization of beer by RO in the 1980s. In the last decade, membrane ﬁltration has established itself for the clariﬁcation of wine, beer and vinegar and based on its now proven reliability in other production steps. 1.3.1 Beer

The conventional brewing process starts in the brew house with the stepping of the malt with hot water to produce wort, a thick sweet liquid. The wort is then passed to the wort boiler in which it is brewed/boiled for up to 2 h followed by clariﬁcation and cooling. The clariﬁed and cooled wort is combined with yeast and passed on to the fermentation tanks in which the yeast converts the grain sugar to alcohol and as such produces beer. Before being transferred to the bright beer tanks, the beer is commonly clariﬁed. The ﬁnished beer might then be ﬁne-ﬁltered and pasteurized before bottling. In the case of beer dealcoholisation, the alcohol removal takes place before the beer clariﬁcation. The overall brewing process with potential applications of cross-ﬂow membrane ﬁltration is shown in Figure 1.5. 1.3.1.1 Beer from Tank Bottoms/Recovery of Surplus Yeast After fermentation, yeast is settling at the bottom of the fermentation vessels. The settled tank bottoms account for 1.5–2% of the total beer volume and, apart from the

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Figure 1.5 Beer production with membrane technology.

yeast, contain a high proportion of beer that is lost if not recovered. In order to recover the beer and concentrate the yeast up to 20% DM, a continuous membrane process has been developed, which separates the beer from the yeast by cross-ﬂow MF with plate-and-frame modules or tubular modules. The layout of this process with plateand-frame modules is shown in Figure 1.6. The investment and operating costs of the beer recovery plant are balanced by the beer recovered from the yeast. For a typical brewery with an annual production of

Figure 1.6 Recovery of beer and surplus yeast from tank bottoms.

1.3 Fermented Food Products

Figure 1.7 Concept of beer clarification by MF.

2 million hl, the recovered beer amounts to 24 000 hl, or about 1% of the annual production [16]. Furthermore, the recovered yeast has an increased dryness that supports further processing. 1.3.1.2 Beer Clarification In the traditional brewing process, the clariﬁcation of the beer after fermentation and maturation is often achieved by a separator followed by kieselguhr ﬁltration, a process that is associated with handling and disposal of the powder as well as large amounts of efﬂuents. To overcome these problems, cross-ﬂow MF with plate-and-frame cassettes has been adopted to remove yeast, micro-organisms and haze without affecting the taste of the beer. The concept of this process is shown in Figure 1.7. 1.3.1.3 Beer Dealcoholization The demand for low-alcohol and alcohol-free drinks has been constantly growing over the last decade. The market development, for example in Germany shows an increase in the annual consumption of alcohol-free drinks from 130.4 l per person in 1980 to 248.4 l per person in 1999, while in the same period the consumption of alcoholic drinks decreased from 179.5 to 156.3 l per person [17]. RO can be used to reduce the alcohol concentration 8–10 times, while maintaining the beer ﬂavor. The dealcoholization of beer by RO is divided into four steps:

Preconcentration – the beer is separated into a permeate stream containing water and alcohol and a retentate stream consisting of concentrated beer and ﬂavours. 2) Diaﬁltration – addition of desalted and deoxygenized water to balance the volume removal with the permeate combined with continuous water and alcohol removal with the permeate. 3) Alcohol adjustment – ﬁne tuning of taste and alcohol content by addition of desalted and deoxygenized water. 4) Post-treatment – to balance taste losses due to removal of the taste carrier alcohol, components such as hops and syrups are added to the dealcoholized beer.

1)

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All the steps are operated at temperatures of 7–8 C or lower, resulting in a highquality beer, the ﬂavor of which is not affected by a heating process. After dealcoholization, the beer is clariﬁed before bottling. 1.3.2 Wine

The traditional wine-making process starts with the crushing and pressing of the grapes followed by must correction, if required. The grape juice from the pressing is centrifuged and transferred to the fermentation tanks, where the fermentation process starts under the addition of yeast. When the fermentation is completed, the yeast fraction from the wine is removed and the wine is moved into barrels for aging. After the aging, the mature wine is clariﬁed, tartar stabilized, sterile ﬁltered and bottled. Membrane processes can replace several of the different separation steps involved in the traditional wine production as shown in Figure 1.8. When the taste of the wine has been deteriorated or dealcoholization of the wine is desired, then these steps are taken before the sterile ﬁltration.

Figure 1.8 Membrane processes in the wine production.

1.3 Fermented Food Products

Figure 1.9 Batch plant for must correction by RO.

1.3.2.1 Must Correction As an alternative to chaptalization or other treatments, RO can be applied to increase sugar contents in the wine without addition of nongrape components at ambient temperatureandtoadjustandbalance thecompositionof themust.Theuseof ROleads to enrichment in tannins and organoleptic components by water reduction between 5 and 20%. This method is particularly suitable to reverse the dilution of the must quality due to rain during the harvest by the selective removal of excess water. However, applying this method to must from grapes of stalled maturity due to cold weather was found to be less effective, since apart from sugar, acid and green tannins are also concentrated [18]. In general, the use of this method is limited by the legislation in the different countries. In Figure 1.9, the concept for a must correction plant is shown. 1.3.2.2 Clarification of Wine The traditional ﬁning after fermentation often involves several steps of centrifugation and kieselguhr ﬁltration to obtain the desired quality. The use of MF/UF can reduce the number of steps by combining clariﬁcation, stabilization and sterile ﬁltration in one continuous operation and eliminates the use of ﬁning substances and ﬁlter material. The key to success in the clariﬁcation of wine is the membrane selection with regard to fouling behavior and pore size. Another important factor is the membrane pore diameter. In Table 1.1, a selection of critical wine compounds and their sizes is given. Typically, MF membranes with pore diameters between 0.20 and 0.45 mm are used for white wine and between 0.45 and 0.65 mm for red wine ﬁltration. 1.3.2.3 Rejuvenation of Old Wine (Lifting) Aging might deteriorate the taste of wine viniﬁed to be consumed young. A diaﬁltration process by RO can be applied to lift the wine by removing the negative

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Table 1.1 Wine compounds and sizes [19–22].

Component

Size

Large suspended solids Yeast Bacteria Polysaccharides Proteins, tannins, polymerized anthocyanins Simple phenols, anthocyanics Ethanol, volatiles

50–200 mm 1–8 mm 0.5–1.0 mm 50 000–200 000 D 10 000–100 000 D 500–2000 D 20–60 D

aroma components causing the stale taste with the permeate. The wine is treated by an RO unit, which concentrates the wine slightly by removing mainly water, little alcohol and the negative aroma components. The volume lost by the permeate may be replaced by continuously adding demineralized water to avoid remineralization of the wine. The diaﬁltration process slightly decreases the alcohol content of the wine but improves the quality of the old wine so that it can be sold at a higher price or blended with younger wine. The advantage of this lifting process is that it does not change the structure and composition of the wine, while the effect of the alcohol reduction is minor. 1.3.2.4 Alcohol Removal Similar to the beer market, the demand for low alcohol wine has increased in recent years. Initial trials in the production of alcohol-free wine can be dated back to 1908 when Jung [23] took out a patent on the thermal dealcoholization of wine. Presently, RO is used to remove ethanol and water, which have a relatively low molecular weight in comparison to the other compounds in wine, see Table 1.1, which passes through the membrane, while the larger compounds of the wine matrix are rejected. The process is similar to the dealcoholization of beer, see Section 1.3.1.3, and can be similarly subdivided in preconcentration, diaﬁltration and alcohol adjustment. Apart from producing alcohol-free wines, this technique can be used to adjust the alcohol level in wine. Wine makers often allow their grapes to ripen until an optimum rich ﬂavor is achieved. At this stage, the grape juice often contains high sugar levels, which result in high alcohol content after fermentation. The alcoholic aroma, however, suppresses other ﬂavors in the wine. By use of RO, the wine can be slightly concentrated by removing water and part of the alcohol. This allows wine makers to harvest grapes depending on the grape ﬂavor ripeness and independent of their sugar contents. 1.3.3 Vinegar

The production of vinegar is an old process, referred to in the history as far back as Babylon 5000 BC. Over the years, the product has been developed according to nationality and tradition, resulting in widely different methods of production.

1.4 Fruit Juices

Figure 1.10 Membrane technology in vinegar production.

Vinegar is produced by an aerobe fermentation of bacteria (genus acetobacter) reacting on dilute solutions of ethyl alcohol such as cider, wine, fermented fruit juice or dilute distilled alcohol. The different raw materials (apples, grapes, malt, rice, etc.) each contribute to giving the vinegar its special aroma and ﬂavor. In the traditional production process, vinegar requires a reaction time between 3 and 6 months for formation and sedimentation. For some vinegar types, ﬁning agents are also necessary, which are added to the vinegar after fermentation. The ﬁnal ﬁltration takes place after storage in order to remove the colloids formed. In Figure 1.10, the production process of vinegar including membrane technology is shown. 1.3.3.1 Clarification of Vinegar The clariﬁcation of vinegar by UF is positioned directly after the fermentation step and can substitute many steps in the traditional production. The vinegar ﬁning by UF can be applied for a wide range of vinegar types and results in a vinegar product on the permeate side, that has similar color and organoleptic qualities to the original vinegar but no turbidity. Additionally, proteins, pectins, yeast, fungi, bacteria and colloids are removed and thus the ﬁltration/sedimentation and the clariﬁcation are substituted and the storage time reduced. Hence, the permeate from the UF step can be directly pasteurized before bottling or additional processing. However, UF cannot give the vinegar the aroma, which is normally obtained during storage. This aroma is secured by the storage time in the wholesale and retail stages instead.

1.4 Fruit Juices

The general production ﬂow in the fruit juice industry starts with grinding or crushing of the fruits into an optimal and uniform size of particles and then pressing

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Figure 1.11 Membrane processes in fruit juice production.

out the fruit mash. The traditional ﬁning process consists of long retention time in tanks followed by kieselguhr ﬁltration and requires large amounts of enzymes, gelatin and other chemicals. After clariﬁcation/ﬁning, the fruit juice is concentrated to reduce costs for transportation and storage. The common approach to concentrate fruit juice is by using an evaporator combined with an aroma-recovery unit concentrating the apple juice from originally 11–12 Brix to over 70 Brix. The concentrated fruit juice can then be optionally pasteurized before transportation. The general fruit juice production process including membrane processes is shown in Figure 1.11. 1.4.1 Fruit-Juice Clarification

The clariﬁcation of fruit juice, mainly apple but also grape, pineapple and orange juice by UF has proven to be an attractive substitute for the traditional ﬁning and ﬁltering process from an economic and qualitative point of view since the 1970s. The UF process removes the suspended solids and other high molecular solids and the ﬁltered juice obtains a clarity and excellent quality, which has not previously been obtainable. Thus, the UF process substitutes the ﬁning step in the traditional process. In order to achieve high yield, high capacity and excellent quality, an enzyme treatment and proper preﬁltration must be carried out before the UF system is utilized. Until now, the industrial standard is to use polymeric and

1.5 Other Membrane Applications in the Food Industry

Figure 1.12 Juice clarification (left) and juice concentration (right).

ceramic tubular modules for the clariﬁcation of the juice. However, this module type is associated with low packing density and high membrane replacement costs. Furthermore, this process is commonly run in batch mode and diaﬁltration water has to be added in the ﬁnal stage of the clariﬁcation to maximize the process yield. More recently, a new concept has been developed, which combines a high-speed separator with spiral-wound UF modules to overcome these limitations [17], see Figure 1.12. 1.4.2 Fruit-Juice Concentration

For the concentration of apple juice, the combination of RO and evaporation can provide an interesting process combination. RO as initial step can remove more than 50% of the water content prior to evaporation, while maintaining 98–99% of sugar and acid as well as 80–90% of volatile ﬂavours in the concentrate, see Figure 1.12. By applying RO, concentration levels of 20–25 Brix can be achieved, while the subsequent evaporation can boost these levels to above 75 Brix. By applying this concept, only 7–9 kWh per m3 fruit juice are required, which represents an energy saving of 60–75% compared to direct evaporation. Furthermore, the permeate from the RO unit can be recycled as process water.

1.5 Other Membrane Applications in the Food Industry

Apart from the production processes discussed above there are many other applications of membrane processes in the food industry. The ﬁrst part of this

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section provides an overview of other key membrane applications in the food industry directly related to the product stream. The aim is not to give a complete listing of all possible applications but to document the diverse applicability of membranes in the food production. The second part of this section focuses on the membrane applications in the food industry related to process water and wastewater. 1.5.1 Membrane Processes as Production Step

The continuous improvement and proven use of membranes in the industry has established membrane technology as a molecular separation unit in a wide range of applications in the food industry. In Table 1.2, a selection of other established membrane applications in the food industry from the continuously growing list of applications is presented. 1.5.2 Membrane Processes for Water and Wastewater

The food industry is one of the largest water-using industries. In the industry, water is used as an ingredient, for initial and intermediate cleaning of the product, and as a key agent in the sanitation of the plant. Depending on the purpose, the requirements for the water vary signiﬁcantly. The water used in the food industry can be generally classiﬁed into three types: Process water – potable water used as an ingredient, is part of or in direct contact with the food. 2) Boiler and cooling water – soft water to avoid scaling and fouling of the cooling and heating equipment. 3) General purpose water – potable, often chlorinated water to rinse raw materials, prepared products, and equipment. 1)

After usage, the different water streams have to be treated as for recycling or for discharge. Membrane processes play an important role in both the pretreatment of the water before usage and post-treatment of the water before recycling or discharge. In Table 1.3, some applications of membranes in the pretreatment and posttreatment of water are summarized.

1.6 Future Trends

It is predicted that membrane processes will continue to grow at average annual growth rates of 5–8% in the foreseeable future. Apart from the worldwide acceptance and use of membrane processes, the key drivers for this development can be related to three key areas, which will be discussed below.

Table 1.2

Selection of other membrane applications in the food industry.

Production step

Animal blood plasma Concentration and puriﬁcation of blood plasma

Membrane processes

Comments

UF

Concentration up to 30% total solids (TS). Low molecular weight components are removed with permeate, for example, salts. Diaﬁltration can increase purity. Concentration of high molecular weight peptides in retentate.

NF/RO

Volume reduction before spray drying.

Egg Whole-egg concentration

UF

Concentration up to 40–44% TS. Low molecular weight components are removed with permeate, for example, salts and sugars.

Egg-white concentration

UF

Concentration up to 20–21% TS. Puriﬁcation by removing salts, glucose and other low molecular components with permeate.

RO

Concentration up to approx. 24% TS. Product loss less than 0.05% of the solids in the feed.

Gelatin and gums Agar and agarose concentration

UF

Concentrate up to 2% TS (agarose) and 4–5% TS (agar). Removes more than 50% of water.

Carrageenan concentration

UF

Concentration up to 3–4% carrageenan. Puriﬁcation and decolorization by removing low molecular carrageenan, salt, color and sugars.

Apple and citrus pectin concentration

UF

Concentration up to 4–7%. Puriﬁcation by removing low molecular components, for example, salt and sugars.

Gelatin concentration

UF

Concentration of gelatin up to 25% depending on grade of hydrolytic conversion and bloom value.

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UF

Concentration of blood cell fraction

1.6 Future Trends

Recovery of peptides from blood-cell fraction

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Table 1.3 Process and wastewater.

Production step

Membrane processes

Comments

Water pre-treatment Desalination/softening of process, boiler and cooling

NF/RO

RO removes minerals, particles plus most of the bacteria and pyrogens.

Preparation of diaﬁltration water

RO

Diaﬁltration water is high-quality water in accordance with process water standards.

Pyrogen removal

UF, NF, RO

Membranes with MWCO less than 10 000 remove most pyrogen.

Water post-treatment Concentration of sugar water

RO

Concentration of sugars to reduce BOD. Water and sugars might be recycled in the process.

Concentration of food proteins

UF

Concentrated food proteins, for example from the washing step can be concentrated and reused.

Condensate polisher

UF, NF, RO

Concentration of the evaporator condensate, for example in case of carry-over with high BOD/COD.

Concentration of UF permeate

RO

UF permeate contains the low molecular components such as sugars and salts.

Biological treatment

MF/UF

Membrane bioreactor (MBR) with water removal by MF/UF.

1.6.1 New Applications of Membrane Processes

The development of new applications of the established membrane processes MF, UF, NF and RO will be driven by economical and environmental targets. An additional driver for membrane processes is the high growth rate of the market for functional foods, a segment in which membranes has a high potential. In Table 1.4, some of the most recent research trends on membrane applications for MF, UF, NF and RO in the food industry are summarized. 1.6.2 New Membrane Processes

In recent years, three new membrane processes have been developed for applications in the food industry. The processes and their potential in the food industry are shown in the following.

1.6 Future Trends Table 1.4 New applications of MF, UF, NF and RO in the food industry [9, 24–26].

Application

Membrane processes

Dairy Concentration of whole and skim milk Partly demineralized WPC (baby food, special WPC products) Production of whey protein concentrates and isolates Defatting of whey for high protein WPC Standardization of the protein content in cheese milk

RO NF UF MF MF

Wine Preclariﬁcation of grape juice

MF/UF

Fruit juices Clariﬁcation of pulpy tropical fruit juices Concentration of tomato juice

MF MF and RO

Other applications Concentration of chicken blood plasma Filtration of extra virgin olive oil Dry degumming of vegetable oil

MF/UF UF/NF

1.6.2.1 Pervaporation While the use of pervaporation for the dehydration of organic compounds is stateof-the-art in the industry, the use of pervaporation for the recovery of organic compounds from aqueous solutions is still limited. The key features of pervaporation are the mass transfer of components through a commonly non-porous polymeric or zeolite membrane combined with a phase change from liquid to vapor. The driving force of pervaporation is an activity difference between the feed and permeate side, while the mass transfer can be described based on the solution diffusion model. For the food industry, three potential applications have been under investigation:

Removal of alcohol from wine – a concept has been patented by Lee et al. [27] by using hydrophilic membranes and is carried out similarly to alcohol removal by RO. 2) Aroma recovery from raw material (fruit juices, beer, herbal and ﬂowery extracts) – a commercial process has been developed and successfully tested at a fruitjuice concentrate company [28]. 3) Recovery of aroma components during fermentation – pilot-scale experiments during the fermentation of wine demonstrated the feasibility to recover the complex wine aroma [29].

1)

Pervaporation is, however, despite its successes and potentials, so far not established in the food industry. 1.6.2.2 Electrodialysis Electrodialysis is used to separate uncharged molecules from charged molecules and is therefore used for, for example, the separation of salts, acids, and bases from

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aqueous solutions. The key advantage over other membrane processes is the selectivity of electrodialysis towards charged molecules without affecting uncharged molecules. The driving force of the process is based on a gradient of the electrical potential and the separation is achieved based on the Donnan exclusion mechanism using ion-exchange membranes. This mechanism enables electrodialysis to enrich and concentrate electrically charged ions from aqueous solutions. Potential applications in the food industry are, for example: 1)

Tartaric stabilization of wine by removing potassium, calcium cations and tartrate anions – has been commercialized and is recognized by the International Wine ofﬁce as good practices [30]. 2) Lactic-acid recovery from fermentation broth – realized on a commercial scale to improve productivity. 3) Whey demineralization – effective demineralization after concentration by NF, used in the dairy industry. The use of electrodialysis in some applications is well established in the food industry but the market share of electrodialysis is small compared to MF, UF, NF and RO. 1.6.2.3 Membrane Contactors – Osmotic Distillation The concept of membrane contactors was developed during the 1970s, however, the commercialization of the Celgard Liqui-CelÒ hollow-ﬁber module in 1993 led to the breakthrough of this technology. Membrane contactors are devices that achieve a gas/ liquid or liquid/liquid mass transfer of one phase to another without dispersion by passing phases on both sides of a microporous membrane. Controlling the pressure difference between the two phases carefully, one of the phases can be immobilized in the pores of the membranes and an interface between the two phases can be established at the mouth of each pore. The driving force of the process is the concentration and/or pressure difference between the feed and the permeate side and mass transfer is based on distribution coefﬁcients. Selected applications in the food industry are:

1) 2) 3) 4) 5)

Bubble-free carbonation of soft-drinks – realized in the Pepsi bottling plant in West Virginia to carbonize about 424 l of beverage per minute. CO2 removal followed by nitrogenatation – used in the beer production to preserve the beer and to obtain a dense foam head. Deoxygenized water – water for the dilution of high-gravity brewed beer [31]. Alcohol removal by osmotic distillation – has been tested for wine but not commercialized. Concentration of fruit juices by osmotic distillation – achieves concentrations greater than 60 Brix.

Membrane contactors are currently one of the most active ﬁelds of membrane process and application development with many interesting spin-offs for the food industry.

References

1.6.3 Integrated Process Solutions: Synergies and Hybrid Processes

The development of integrated process solutions such as synergies and hybrid processes is one relatively unexplored area of process development. Until now, commonly only one unit of operation is considered to achieve a predeﬁned separation. Combinations of conventional processes such as centrifugation, evaporation, liquid–liquid extraction and adsorption with membrane processes are rarely used, even though they might offer economical beneﬁts to the end user. However, by integrating membrane processes in their product range, more and more system builders combine the conventional processes with membrane technology. Hence, it seems reasonable to assume that the economic beneﬁts of such process combinations and a wider understanding within the industry of their potentials will support the long-term growth of membrane technology. Overall, cross-ﬂow membrane processes have established themselves in the food industry and many exciting developments will ensure their importance for the future.

Acknowledgements

The author would like to thank Dr Olga Santos for her valuable input to the section on the dairy industry and Prof. Gun Tr€agardh for the critical review of the manuscript. Further, the author would like to thank Hanne Jonck for proofreading the manuscript.

References 1 Holm, S., Malmberg, R., and Svensson, K.

2

3

4

5

(1986) Method and plant for producing milk with low bacterial content. Patent WO86/001687. Meersohn, M. (1989) Nitrate-free cheese making with the Bactocatch. North Eur. Food Dairy J., 55, 108–113. Hansen, R. (1988) Better market milk, better cheese milk, better low heat milk powder with Bactochatch treated milk. North Eur. Food Dairy J., 54, 39–41. Ottosen, N. and Kønigsfeldt, P. (1999) Microﬁltration of cheese brine. Preliminary, APV Nordic, Membrane Filtration, Silkeborg; Denmark. Quiñones, H.J., Barbano, D.M., and Phillips, L.G. (1997) Inﬂuence of protein standardization by ultraﬁltration on the viscosity, colour, and sensory properties of

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skim and 1% milk. J. Dairy Sci., 80, 3142–3151. Rosenberg, M. (1995) Current and future applications of membrane processes in the dairy industry. Trends Food Sci. Tech., 6, 12–19. Merin, U., Gordin, S., and Tanny, G.B. (1983) Microﬁltration of sweet cheese whey. New Zealand J. Dairy Sci. Tech., 18, 153–160. Fauquant, J., Vieco, E., and Maubois, J.-L. (1985) Clariﬁcation des lactoserums foux par agregation ythermocalcique de la matiere grasse residualle. Lait, 65, 1–20. Maubois, J.-L. (1997) Current uses and future perspectives of MF Technology in the dairy industry. In: Bulletin of the International Federation No. 320, International Dairy Federation, Brussels.

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10 Jost, R. (1997) Cross-ﬂow microﬁltration-

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An extension of membrane processing of milk and whey. In: Bulletin of the IDF 320, International Dairy Federation, Brussels. Karleskind, D., Laye, I., and Morr, C.V. (1995) Chemical pre-treatment and microﬁltration for making delipized whey protein concentrates. J. Food Sci., 60, 221–226. Maubois, J.-L. and Ollivier, P. (1992) Milk protein fractionation, in New Applications of Membrane Process, IDF Special Issue No. 9201, International Dairy Federation, Brussels, Belgium. Tamime, A.Y. (1993) Modern cheese making: Hard chesses, in Modern Dairy Technology (ed. R.K. Robinson), Elsevier Applied Science LDT, New York. Qvist, K.B. (1987) Objective and sensory assessment of texture of danbo cheese made from milk concentrated 2-fold using ultraﬁltration. 272, Beretning/ Statens Mejeriforsøg, Hillerød, Denmark. Qvist, K.B., Thomsen, D., and Kjærgard, J.G. (1986) Fremstilling af havarti ost ud fra mælk, der er koncentreret ca. 5 gange ved ultraﬁltrering. 268, Beretning/Statens Mejeriforsøg, Hillerød, Denmark. Lipnizki, F. (2005) Optimisation and integration of membrane processes in the beverage industry, 10. Aachener Membran Kolloquium, 16–17.03.2005, Aachen, Germany. Gebhardt, W. (2001) Weltforum der Weinund Saftbereitung. F&S Filtrieren und Separieren, 15, 234. Smith, C. (2002) Applications of reverse osmosis in winemaking, wwwvinovation.com. Millet, V. and Lonvaud-Funel, A. (2000) The viable but non-culturable state of wine micro-organisms during storage. Lett. Appl. Microbiol., 30, 136. Peri, C., Riva, M., and Decio, P. (1988) Cross-ﬂow membrane ﬁltration of wines: Comparison of performance of ultraﬁltration, microﬁltration and

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intermediate cut-offs membranes. Am. J. Enol. Vitic., 39, 162–168. Toland, T.M., Fugelsang, K.C., and Muller, C.J. (2009) Methods for estimating protein instability in white wine: A comparison. Am. J. Enol. Vitic., 47, 111. Boulton, R.B., Singleton, V.L., Bisson, L.F., and Kunkee, R.E. (1996) Principles and Practices of Winemaking, Chapman & Hall, New York, NY. Jung, C. (1908) Verfahren, um aus Fl€ ussigkeiten, die ﬂ€ uchtige Riechstoffe und Alkohole enthalten, durch Distillation den Alkohol und die Riechstoffe getrennt zu gewinnen, Swiss Patent 44,090 Torres, M.R., Marın, F.R., Ramos, A.J., and Soriano, E. (2002) Study of operating conditions in concentration of chicken blood plasma proteins by ultraﬁltration. J. Food Eng., 54, 215–219. Vaillant, F., Millan, A., Dornier, M., Decloux, M., and Reynes, M. (2001) Strategy for economical optimisation of the clariﬁcation of pulpy fruit juices using cross-ﬂow microﬁltration. J. Food Eng., 48, 83–90. Bottino, A., Capannelli, G., Comite, A., Ferrari, F., Marotta, F., Mattei, A., and Turchini, A. (2004) Application of membrane processes for the ﬁltration of extra virgin olive oil. J. Food Eng., 65, 303–309. Lee, E.K., Kalyani, V.J., and Matson, S.L. (1991) Process for treating alcoholic beverages by vapor-arbitrated pervaporation, US Patent 5,013,447 Partos, L.(30th Nov. 2004) http:/ foodproductiondaily.com/news/ printNewsBis.asp?id¼56425. Sch€afer, T., Bengtson, G., Pingel, H., B€ oddeker, K.W., and Crespo, J.P.S.G. (1999) Recovery of aroma compounds from wine-must fermentation by organophilic pervaporation. Biotechn. Bioeng., 62, 412. Eurodia Industries (2002) Tartaric stabilisation of wine, www.eurodia.com. Gableman, A. and Hwang, S.-T. (1999) Hollow ﬁbre membrane contactors. J. Membrane Sci., 159, 61.

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2 Membrane Processes for Dairy Fractionation Karin Schro€en, Anna M.C. van Dinther, Solomon Bogale, Martijntje Vollebregt, Gerben Brans, and Remko M. Boom 2.1 Introduction

Traditionally, milk has been separated in order to produce a wide range of dairy products. In some cases, separation is minimal, such as in regular full-fat milk, which is standardized in order to have the correct amount of fat. But for semi-skim milk and skim milk, separation needs to be done more rigorously because a large amount of the milk fat or even all of it needs to be removed in order to obtain the desired fat percentage in the product. When considering complex dairy products, such as cheese, it is clear that not the entire milk is used but only partly, and valuable byproducts are generated (see Table 2.1). From the milk, the cream and casein fraction (main component of cheese) are separated, after which a certain amount of fat is added back according to speciﬁcations for the type of cheese that is to be prepared. For Gouda cheese production, rennet, CaCl2, and starter culture are added, after which a gel is formed by the casein, which is subsequently cut to small pieces (curd) to remove the so-called whey. The curd particles are subsequently pressed into cheese shape, brined, and stored. The whey contains the so-called whey proteins, and these are of considerable value, since they are easily digestible, and are added to, for example, sport drinks. For the separation of fat from milk, mostly centrifuges are used, but membranes could be an interesting alternative, which is explained in Section 2.2.1. Besides separation of milk fat, also all other milk components are in a range in which membranes are effective (for a summary see Table 2.2). In Figure 2.1, a general comparison is made between the size of the dairy components and the pore size of membranes. In some ﬁelds, membranes have established their value such as processing of whey and they are gaining popularity in other dairy applications as described in Dauﬁn et al. [1]; for a recent review paper see Brans et al. [2]. However, separation of milk in many different fractions has not been described in the literature; mostly papers focus on a single stage. Some successful examples are the separation and fractionation of fat globules, the reduction of bacteria and spores in skim milk, concentration of casein micelles (for cheese manufacturing), and puriﬁcation of serum proteins, and

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Table 2.1 General overview of processing steps required in Gouda cheese preparation including some by-products (reprinted from Brans et al. [2] with permission from Elsevier).

Processing step/Separation

Product

By-product

Separation of cream from rest and standardization fat content Curding through addition of rennet, CaCl2, and starter Cutting of curd Pressing of curd particles Brining Ripening

Cheese milk

Cream or skim milk depending on fat content

Gelled milk Curd Shaped cheese Salted cheese Mature cheese

Whey Brine with cheese components

these will be discussed in the next section. In general, it can be mentioned that the various membrane processes that are discussed in the literature for dairy applications have a number of aspects in common related to ﬂux decline and fouling, and related to that selectivity. Logically, many papers deal with strategies to prevent ﬂux and selectivity reduction, for example, the uniform transmembrane pressure is developed to have similar conditions over the entire length of the membrane, and these strategies will be discussed in Section 2.3. Since ﬂux and selectivity loss also originate from the membrane speciﬁcations; a pore-size distribution is expected to inﬂuence the sharpness of the separation, we will also discuss membranes with more uniform pore sizes such as asymmetric ceramic membranes, track-etched membranes [3], silicon microsieves [4], and metal microﬁlters [5]. Table 2.2 Average composition of cow milk: concentration and size distribution (reprinted from Brans et al. [2] with permission from Elsevier).

Concentration in whole milk (g/l) Water Fat globules Casein (in micelles) Serum proteins a-lactalbumin b-lactoglobulin BSA Proteose-pepton Immunoglobulins Lactoferrin Transferrin Others Lactose Mineral substances Organic acids Other

87.1 4.0 2.6 0.7 0.12 0.32 0.04 0.08 0.08 0.01 0.01 0.04 4.6 0.7 0.17 0.15

Size range and average (at weight average) 0.1–15 mm, average 3.4 mm 20–300 nm, average 110 nm 3–6 nm 14 kD 18 kD 66 kD 4–40 kD 150–900 kD 86 kD 76 kD 0.35 kD

2.2 Membrane Separation of Components

Somatic cells 10 m

Fat globules 1 m

Bacteria and spores MF

100 nm

Casein micelles Casein submicelles

10 nm

Serum proteins 1 nm

Lactose Salts

0.1nm

UF

Water

NF RO

Figure 2.1 Comparison of size of components in milk and pore size of membranes. MF: microfiltration; UF: ultrafiltration; NF: nanofiltration; RO: reverse osmosis (reprinted from Brans et al. [2] with permission from Elsevier).

2.2 Membrane Separation of Components 2.2.1 Removal of Milk Fat from Whole Milk

As mentioned previously, mostly centrifugation is used for separation of milk fat from milk, although membrane separation is technically possible, as indicated in a patent by Alfa-Laval [6]. An advantage of using membranes instead of centrifuges could be that the fat globules are less damaged, which is expected to enhance cream stability, and sensory perception. Milk-fat droplets range in size from 0.1 to 15 mm, with an average around 3.4 mm. At room temperature, the fat is mostly solid, and in order to avoid clumping the liquid needs to be heated up to 50 C. Goudedranche et al. [7] who were mainly interested in consumer perception of cream, describe the fractionation of milk-fat globules with a 2-mm ceramic membrane, but unfortunately did not report the size distribution of the two fractions. The consumers preferred the small fat globules that gave products with ﬁner texture, to the larger fat globules and a standard cream. Clearly, fractionation of fat particles can lead to products that are more appreciated by consumers, and this should drive further development if only for the cream in milk. 2.2.2 Removal of Bacteria and Spores from Skim Milk (Cold Pasteurization)

The main advantage of using microﬁltration for the reduction of bacteria and spores from milk is that the taste of the milk is not affected because no heat treatment is required. Besides, the reduction that can be achieved is higher than for centrifugation [8], and as a result, the shelf life of milk is extended. Further, microﬁltration has

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Table 2.3 Comparison of cold sterilization results from various sources.

Membrane type and flux

Process conditions cross-flow/pressure, UTP, backpulsing

Log reduction

Source

Ceramic 1.4 mm; 1.4 104 m/s Reversed asymmetric 0.87 mm; 1.4 104 m/s Microsieve 0.5 mm

50 kPa, 7.2 m/s UTP

above 3.5

0.5–1 m/s; backpulsing 0.2–1 s1 dead-end ﬁltration of spiked SMUF 6 to 8 m/s

between 4 and 5

Saboya and Maubois [10] Guerra et al. [11]

Bactocatch: ceramic membranes

6.6

van Rijn and Kromkamp [12] Holm et al. [13]

been described in patent literature as a pretreatment method for skim milk to be used in the production of raw milk cheeses [9]. Various authors have worked on this topic, and they have used rather different methods, and operational conditions. In Table 2.3, the details are summarized, and it is clear that various membranes and conditions have been used, although not all information is displayed in patent literature. In general, the log reductions that can be obtained (10 000 fold reduction or more) are very interesting, and this makes microﬁltration an interesting option for cold sterilization, albeit the log reduction is not as high as obtained by regular heat treatment. The highest log reduction (6.6: higher than for regular pasteurization) was claimed for microsieves, which are silicon plates with very accurately manufactured pores using laser interference lithography. Although the bacterial reduction was measured for deadend ﬁltration of SMUF (simulated ultraﬁltrate) spiked with Bacillus subtilis, over a 0.5-mm microsieve, we believe that the high reduction obtained with this model system for milk is a result of the extremely narrow pore-size distribution of the microsieve. In Figure 2.2, micrographs of a ceramic membrane and a microsieve are shown. 2.2.3 Concentration of Casein Micelles in Skim Milk

As mentioned in the introduction, in cheese production, various waste streams are created, and especially whey is a big waste stream; from 10 l of milk, 1 kg cheese is produced, and therewith also 9 l of whey. Because of these huge volumes that are involved, it is an interesting notion to start cheese making with a concentrated casein solution, and to remove whey proteins and other low molecular weight components. Although casein is only 2.6% weight percentage of milk, it contains a lot of water and is very voluminous. Typical diameters of casein are between 20–300 nm, with an average of 110 nm [14].

2.2 Membrane Separation of Components

Figure 2.2 Micrographs of a microsieve (image courtesy of Aquamarijn) and ceramic membrane.

This topic has attracted the attention of various authors, who used various process conditions and membranes; an overview is given in Table 2.4. Although the studied conditions were rather different, the results were not, maybe with the exception of the work of Krstic et al. [15], who used turbulence promoters. For concentration of casein micelles, control of the membrane ﬂux through control of fouling seems most important, and the fact that some whey protein may end up with the casein and vice versa, is not such an issue. Casein concentration through microﬁltration is a better option compared to the use of traditional ultraﬁltration as pretreatment for cheese (which concentrates both casein and whey protein), since this leads to less whey protein in the cheese process. When comparing casein concentration to cold sterilization, it is immediately clear that separation of bacteria needs to be and remain sharp, and therefore, this separation needs to meet higher demands regarding selectivity than casein concentration, although the economics of the process are affected by the selectivity of the process [16].

Table 2.4 Comparison of casein concentration from various sources.

Membrane type and flux

Process conditions cross-flow/pressure

Ceraﬂo 0.22 mm; 2.5 105 m/s Membralox 0.2 mm 1.9 105 m/s 1.3 105 m/s Ceramem asymmetric 0.05 mm; 3.1 105 m/s Membralox 0.1 mm; 9.7 105 m/s 2.5 104 m/s

6.9 m/s; 190 kPa

3

Pouliot et al. [17]

7.2 m/s; 193 kPa

2

Vadi and Rizvi [18]

5.4 m/s; 138 kPa

10 2

0.45 m/s; 34 kPa turbulence promoters 12.5 m/s; 65 kPa ( þ TP)

Concentration factor

1

Source

Punidadas and Rizvi [19] Krstic et al. [15]

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2.2.4 Recovery of Serum Proteins from Cheese Whey

Traditionally, whey was considered a waste product of cheese making, but nowadays, whey proteins are a considerable source of income for dairy companies. Not surprisingly, separation technology, including membrane separation was developed to capture these valuable components. Whey is mostly high in salt, and therefore, demineralization is needed, and for this electrodialysis or ion-exchange resins are used [20], but also nanoﬁltration has been proposed by van der Horst and coworkers [21]. An added beneﬁt of nanoﬁltration is that it reduces energy consumption, and the partially demineralization product can be spray dried and used in food or feed applications. In the work of Doyen and coworkers [22], various membranes were compared among which were polymeric (PSF/PVP), ceramic (ZrO2) and organomineral (ZrO2/PSf) membranes, and they found that the plateau ﬂuxes were comparable; the fouling layer was the limiting factor in whey protein concentration and not the permeability of the membrane. Since all proteins are retained, prevention of gel formation is critical for process operation. Various proteins are present is whey, which are all of considerable economic worth, such as a-lactalbumin, b-lactoglobulin, bovine serum albumin, immunoglobulins, lactoferrin, transferrin, and some minor proteins and peptides (see also Table 2.2). For example, b-lactoglobulin can be used in emulsiﬁcation, foaming and gelling [23], and for lactoferrin and a-lactalbumin there are pharmaceutical applications [1, 24]. Further, there is an increasing interest in bioactive hydrolysates from serum proteins [25]. The reported separation methods for these proteins include thermal aggregation of a-lactalbumin [26], ion-exchange chromatography, precipitation, ultraﬁltration or a combination of these methods [27–31]. Besides, it was shown to be possible to enhance the selectivity of an ultraﬁltration process by adjusting pH and salt to inﬂuence electrostatic and steric interaction [23, 32]. From the previous sections, it is clear that various separations such as fat separation, cold sterilization, casein concentration, and whey-protein isolation, have been carried out successfully using membranes. However, one factor limits milk fractionation and this is ﬂux decrease related to fouling. Design parameters that can be used to control this are discussed in the next section.

2.3 Methods to Enhance Membrane Separation

As mentioned in the previous section, the accumulated layer or fouling layer determines membrane behavior in many dairy separations. It is generally accepted that it is not the membrane but (the rate of) accumulation that is the limiting factor for membrane ﬁltration of milk [33], although different authors point to different aspects of the accumulated layer as being most relevant [34, 35]. This is also due to the various methods that have been used to access the fouling layer such as SEM (e.g., [34]), AFM (e.g., [35]), ATR–FTIR and EDX [36], streaming-potential measurements [37], and

2.3 Methods to Enhance Membrane Separation

ﬂux measurement, in combination with retention measurement as is regularly used (e.g., [38]). An overview of the various methods used to assess membrane fouling can be found in a recent review by Le-Clech et al. [39], and the relation between membrane surface morphology and membrane performance is described comprehensively by Khulbe et al. [40]. For simplicity reasons, in this section we will use the term ﬂux decrease for any effect that causes this instead of fouling. Flux decrease may thus be linked to concentration polarization, cake ﬁltration, adsorption, depth fouling, pore blocking, or any other effect that reduces the ﬂux. In spite of the different interpretations of membrane fouling/accumulation of components, a number of concepts have been developed to keep the ﬂux at acceptable levels, and these will be discussed ﬁrst. To limit ourselves, we will discuss methods that act on short-term ﬂux decrease, and will not discuss cleaning methods, which are needed to mediate long-term ﬂux decrease, and codetermines the lifetime of a membrane. In section 2.5, we will discuss particle and component behavior in more detail, in relation to speciﬁc aspects of ﬂux decrease, and show how this can be used to design separation processes. 2.3.1 Critical Flux Concept

Transmembrane flux (m/s)

In the critical ﬂux concept proposed by Field and coworkers [41, 42] and recently reviewed by Pollice [43] for membrane bioreactors, three regions are distinguished, as schematically indicated in Figure 2.3. In region I, the transmembrane pressure is below the critical pressure and the ﬂux is linearly dependent on the applied pressure. This dependency can be determined by the clean-water ﬂux as stated in the hard form or lower than the clean-water ﬂux, which is the weak form of the critical ﬂux criterion. Filtration in this region is also known as subcritical ﬂux operation and is advised to obtain optimal selectivity, since accumulation is minimal, due to the low applied

II

I

III III Jcrit

Pcrit Trans membrane pressure (Pa) Figure 2.3 Schematic representation of the critical flux concept. In region I, the flux is linearly dependent on pressure until at a critical pressure (Pcrit) the critical flux (Jcrit) is reached.

The flux levels of as a function of pressure in region II, and even decreases in region III when the pressure is increased further (reprinted from Brans et al. [77] with permission from Elsevier).

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pressures. Because of the low pressure, the ﬂux values are low and the required membrane area necessarily high. In region II, the ﬂux is no longer linearly dependent on the transmembrane pressure, and the ﬂux may be determined by the accumulated layer. The value of the ﬂux can be estimated with gel ﬁltration model and/or backtransport models (e.g., [44]). Although selectivity of the membranes may be inﬂuenced in this region, it is still often chosen because it allows best use of the installed surface area when considering only volumetric productivity, regardless of selectivity. In region III, the applied pressure it too high to maintain an acceptable ﬂux, and mostly this is related to cake formation and compaction. If a membrane process is to be operated in region III, it is necessary to remove the deposited layer at short intervals, for example, through frequent backpulsing. When considering the dairy processes presented in the previous section, in relation to the critical ﬂux concept, it should be mentioned that reduction of bacteria and spores, and concentration of casein micelles is carried out near the critical pressure. Concentration of whey protein is carried out in region II in order to minimize the membrane area, while isolation of whey proteins has to take place in region I for selectivity reasons. In all regions, adsorption of components to the membrane surface can take place, and this can lead to ﬂux loss, and related to this loss of selectivity. In order to prevent this, membrane modiﬁcation may be needed, and this will be presented in a later section, ﬁrst we focus on other processing methods that help keep the ﬂux at acceptable levels. 2.3.2 Uniform Low Transmembrane Pressure Concept (UTP)

In order to increase turbulence in membrane modules, increasing cross-ﬂow velocity is a straightforward option. However, this also results in a pressure gradient across the membrane module, leading to different ﬁltration conditions along the length of the membrane. Since this will inevitably inﬂuence local selectivity, a new concept was proposed, the so-called uniform low transmembrane pressure concept (UTP), which allows a constant pressure drop over the length of the membrane module, for example, through applying a cross-ﬂow on the permeate side [10]. Obviously, this extra cross-ﬂow increases the amount of energy needed during operation but in spite of this, UTP is currently the most popular strategy against ﬂux decrease during the ﬁltration of skim milk to retain bacteria and the concentration of casein micelles. Instead of a cross-ﬂow on the permeate side, membranes can also be adjusted as is the case in Isoﬂux and Gradient Porosity membranes [10]. These membranes have a decreasing membrane resistance over the length of the tube, which has the same effect as UTP, but without the need of a cross-ﬂow on the permeate side. 2.3.3 Turbulence Promotion

In the literature, various options to promote turbulence have been proposed such as vibrating modules [45], rotating-disk modules [46, 47], static mixing inserts [15],

2.3 Methods to Enhance Membrane Separation

spacers, turbulence promoters, and inserts, and the use of Dean vortices or microturbulences [48]. Some methods prevent particle deposition through increased shear rates close to the membrane surface, by either vibration, or rotation. Although interesting effects can be realized through vibration, in general it is difﬁcult to use these equipments on a large scale. Regarding rotation, sealing of the equipment to prevent microbial contamination is an issue, and this may make large-scale installation impossible. The static mixing elements have been shown to increase ﬂuxes (see Table 2.4), and are effective turbulence enhancers, although there are some doubts regarding their cleanability, and the creation of so-called dead areas, which are a source for recontamination by microorganisms. Creation of ﬂow instabilities, such as Dean vortices, is an elegant method to locally increase mass transfer, but may not be suited for many membrane conﬁgurations. 2.3.4 Backpulsing and Flow Reversal

Although turbulence promotion may be one of the side effects of backpulsing and ﬂux reversal, we have decided to dedicate a separate section to them given their relevance for membrane separation (i.e., prevention of ﬂux decrease) in practice. Various terms are in use for the temporary reversal of ﬂow through the membrane, such as backpulsing, backwashing, backﬂushing, and backshocking [49, 50], and in all these cases permeate is pressed back into the feed stream. Through this type of reversal of ﬂow, the deposited components are carried away from the membrane and ideally taken away by the cross-ﬂow. The frequency at which ﬂow is reversed can be high (0.2–1.0 s1) as reported by Guerra and coworkers [11]. These authors reported good results for the reduction of bacteria in skim milk with a combination of UTP and backpulsing (see Table 2.3). Besides reversal of ﬂow through the membrane, the feed ﬂow as such can also be used to improve ﬁltration performance, be it through pulsating ﬂow, or even reversal of ﬂow. In this case, rapid velocity changes occur in the cross-ﬂow channel [51, 52]. Pulsating ﬂow is difﬁcult to use at large scale, because the effect of the pulses is dampened. Of the methods mentioned in this section, in general, high-frequency backpulsing is the method of choice in industrial applications possibly in combination with UTP application. 2.3.5 Other Methods

Many other process options that may aid membrane ﬁltration are known from the literature and they are listed in Table 2.5 in order to make this overview complete; as mentioned previously, (chemical) cleaning as such is not taken into account. Air slugs have been used to locally enhance turbulence [53, 54], but unfortunately, they also induce foaming and protein denaturation in dairy applications. Scouring particles have been used for the same purpose, but they are notoriously hard to reuse and cause damage to the membrane and installation [55]. Acoustic waves and

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Table 2.5 Other methods to enhance membrane performance.

Method

Advantages/disadvantages

Source

Air slugs

Hard to control in large membrane systems; foam formation; protein denaturation

Cui and Wright [53] Cui and Taha [54]

Scouring particles

Hard to control in large membrane systems; reuse of particles; damage to system

Noordman et al. [55]

Acoustic or ultrasonic waves and sonication

Protein denaturation; expensive to scale up

Wakeman and Tarleton [56] Duriyabunleng et al. [57] Villamiel and de Jong [58]

Constant or pulsed electric ﬁelds

Suitable for isolation of whey proteins

Visvanathan and Ben Aim [59] Wakeman [60]

sonication cause vibrations and cavitations, which facilitates transport of particles, but at the same time, they induce denaturation of protein [56–58]. Due to these speciﬁc disadvantages, none of these techniques seems to be promising for application in dairy processing. Electric ﬁelds, either constant or pulsed, have been successfully applied in the separation of whey proteins [59, 60], but because pH adjustment is needed this is not expected to be a viable process for separation of other milk components.

2.4 Use of Models for Membrane Separation

Although it is tempting to use an experimental approach to investigate membrane separation, models can in principle facilitate the design of membrane processes more than any experiment can, although we strongly feel that experimentation and validation are always required. Many models are available in the literature for ultraﬁltration and microﬁltration, predicting various aspects of ﬁltration on different scales, but many are related to the behavior of particles, which are idealized components. Some examples of these models can be found in [61–68]. Most probably, the review papers of Belfort and coworkers [44], and Bowen and Jenner [69] are good starting points for those that are not so familiar with models for membrane ﬁltration. Besides, various descriptive models are proposed, but mostly these models are limited to the speciﬁc apparatus, membrane, and liquids/components for which they were derived, and therefore are of limited use. When testing models against experimental data, there is always the challenge to match the idealized situation of the model, which mimics the physical aspects very well, with the not so ideal situation during ﬁltration. For example, numerous components may be present, the membrane may have a pore-size distribution,

2.5 How to Get from Separation to Fractionation

which inﬂuences the separation, and interactions with the membrane may play a role. It is not always necessary to consider all these aspects, but even selection of the most relevant ones may be a difﬁcult task, although some success stories are also known from the literature. For concentration of casein from skim milk, Samuelsson and coworkers [70] used models with different backtransport mechanisms, and they found that shear-induced diffusion described the observed behavior best. Clearly, basic understanding of particle and component behavior contributes to understanding of the relevant phenomena during separation and the separation characteristics (see next section for another example). Further, computer models were found to be very useful to investigate various aspects of module design such as the liquid ﬂow in relation to cake formation [71], but also the effect of inserts and spacers have been evaluated through CFD [72, 73]. When considering what is done in the ﬁeld of modeling, many aspects have been described well, for example, CFD can be used very well in the design of ﬂow-through modules, however, a link between particle behavior, and separation on the module scale is hard to achieve, also because of the completely different scales at which effects take place. Some interesting studies have recently become available in the literature [74], in which particle behavior is linked to behavior during ﬁltration. Concentration polarization and cake layer build-up on microsieves was investigated for particles that are not able to pass the pores at a ﬁxed cross-ﬂow velocity of 0.32 m s1. Illustrative examples of CFD simulation results are shown in Figure 2.4a. At longer ﬁltration times the layer becomes thicker, and eventually the layer becomes this concentrated that cake layer formation takes place. In Figures 2.4b and c, the pressure dependency of the ﬂux is shown. The CFD simulations have generated very detailed information on the local composition in relation to membrane ﬂuxes, and have proven to be of great value in understanding ﬁltration behavior as well as determining those conditions at which selectivity is expected to be least affected, that is, the critical ﬂux/pressure value can be derived from Figure 2.4. Although the situation in the simulation cannot be translated one on one to milk-ﬁltration experiments because of computational limitations, we still learned valuable lessons that guided us in choosing better process conditions.

2.5 How to Get from Separation to Fractionation

In the previous sections, various aspects have been discussed and some of these we ﬁnd extremely relevant to move from separation to fractionation. More speciﬁcally, we will discuss membranes with uniform pore size, extensive computer simulations on particle behavior, and membrane modiﬁcation here, since they may hold the key to fractionation. First, if the pore size is uniform, the selectivity of the separation is expected to be very sharp (although other options are also available as will be explained in the outlook section). Secondly, modeling of particle behavior is essential to obtain a better understanding of backtransport mechanisms, which in turn will determine the selectivity of a separation in relation to process conditions. Since

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Figure 2.4 Illustration of a CFD simulation on concentration polarization and cake-layer formation during microsieve filtration. (a) the effect of transmembrane pressure on layer build-up; (b) the

steady-state flux as a function of transmembrane pressure; and (c) the flux as a function of time (reprinted from [74] with permission from Elsevier).

components will contact a membrane eventually, membrane modiﬁcation targeted at prevention of adsorption or other initial contacts is also expected to be one of the keys to get to fractionation. 2.5.1 Membranes with Uniform Pore Size

Various membranes are known for their uniform pore size, such as Nuclepore membranes that date as far back as 1962 [75], silicon-based microsieves [12], polymeric microsieves [76], but also metal sieves [77]. Aside from the fact that these membranes are ideal candidates for highly selective separation, they are also an ideal research tool, since pore-size distribution does not play a role. Because microsieves can be made with different pore sizes and geometries, they allow investigation of parameters that otherwise would not have been possible. For example, particle release from various pore geometries was investigated through

2.5 How to Get from Separation to Fractionation

computer modeling, and it was found that particles were released most easily from triangular pores, although from a fractionation point of view this design may not be the ideal choice because the pore is only partially blocked. For fractionation, a round pore is the best choice [78] since it is either blocked and does not contribute to the ﬂux, or is fully selective. In another paper, Brans and coworkers [78] showed the importance of the substructure of the microsieve, which limits the operating ﬂux considerably, but can be resolved through a small change in design. 2.5.2 Simulation of Particle Behavior

Component behavior during ﬁltration is very complex, and this is even enhanced by the size distribution of the components. Based on their size, they may or may not be retained by the membrane, or by the accumulating layer, and size will determine which backtransport mechanisms they will be subjected to. In a classic study by Belfort et al. [44], backtransport mechanisms were linked to permeate ﬂuxes and sizes of the components. In general, Brownian diffusion is the dominating transport mechanism for particles below 0.1 mm and inertial lift is the main mechanism for particles above 10 mm. For particles with intermediate size, which are abundantly available in milk, shear-induced diffusion is the main mechanism of backtransport. It is obvious that for a relevant model, information on the resulting diffusion coefﬁcients is needed in order to come to realistic representations for membrane ﬁltration. Especially, for particles of intermediate size, simulation of their behavior is far from trivial, because the interactions between particles and liquid need to be fully resolved; and this is possible in the Lattice-Boltzmann method [79, 80]. For casein micelles and fat globules, there are indications that they can be treated as hard spheres [81], and this facilitates modeling. Kromkamp [82] has used this approach to investigate the shear-induced diffusion behavior of monodisperse and bidisperse suspension, and the resulting diffusion coefﬁcients can be implemented in ﬁltration models such as described in Section 2.4 for microsieve ﬁltration (see Figure 2.4). 2.5.3 Membrane Modification

As indicated in the previous sections, in milk many components are present (notably proteins) that will interact with membrane surfaces, and mostly will do so in an irreversible way unless subjected to rigorous cleaning. Since any irreversible accumulation inﬂuences the selectivity of the separation, prevention of these interactions is a good way to keep selectivity in place, and this is even more relevant for the previously mentioned microsieves with uniform pore size. For these speciﬁc membranes, we have developed the chemistry to modify them at will [83, 84], including protein repellence through covalent attachment of EO6-containing

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components that reduce the adsorbed amount of BSA and ﬁbrinogen below the detection limit [85, 86].

2.6 Outlook

Although uniform pores, modeling, and modiﬁcation are relevant to mature dairy fractionation, we have to stay open for other opportunities, as is nicely illustrated in the work of Kromkamp et al. [87]. In this case, particle segregation and migration was found to play an overruling role in a speciﬁc dairy separation. Milk-fat globules (sizes ranging from 1 to 10 mm), were to be fractionated with a tubular, ceramic MF membrane with 5.0 mm average pore size, and the transmembrane pressure over the membrane was varied, to keep the permeate ﬂux constant without allowing particle accumulation. In Figure 2.5, the particles size and the relative fat content of the permeate are shown as a function of the applied pressure. For the highest cross-ﬂow velocities, at which particle migration is promoted most, the particle size and fat content are relatively constant, but much lower than in the feed. For the lower crossﬂow velocity, at which particle migration is less pronounced, the particle size and fat content clearly increase with higher ﬂux, while the particle size and fat content almost reach the value in the feed solution at the highest ﬂux measured. Note that these effects cannot be a result of components accumulation since that was excluded in the measurement. This has lead to the conclusion that inside the feed stream segregation (particle migration) has taken place with the larger particles located in the middle of the feed channel, as is depicted schematically in Figure 2.6, and this implies that there is a completely new angle on fractionation, namely through control of the applied

Figure 2.5 Relative fat content and particle size of milk fat globules as a function of the applied transmembrane flux, and cross-flow velocity (reprinted from [87] with permission from Elsevier).

References

Figure 2.6 Schematic representation of migration effects that facilitate membrane fractionation [88].

ﬂux. In this case, the pore size of the membrane is no longer relevant, but simulations of particle behavior and membrane modiﬁcation are still very relevant to make best use of this ﬁnding.

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3 Milk and Dairy Effluents Processing: Comparison of Cross-Flow and Dynamic Filtrations Michel Y. Jaffrin, Valentina S. Espina, and Matthieu Frappart 3.1 Introduction

The dairy industry is very important in Europe where it represents 14% of agricultural national production [1]. The European dairy industry is famous for the quality of its products, especially for its variety of cheeses and yogurts, dairy cream, ice creams, and so on. Milk is a complex ﬂuid and an important source of proteins. The average composition of milk is given in Table 3.1 [2]. As noted by Brans et al. [3], the functionality of milk proteins is larger if they have been separated and puriﬁed. Thus, their fractionment leads to more efﬁcient and diversiﬁed applications. 3.1.1 Properties and Applications of Various Proteins 3.1.1.1 Caseins This protein (24 kDa in molecular size) is generally aggregated as micelles, which average 110 nm in size (or about 300 kDa). Several casein species exist, a1, a2, b, k. Concentrated casein solutions can be mixed with cream for production of cheese and for standardization of milk composition, required for industrial cheese production (between 36 and 45 g/L). They are also used for infant formula and as emulsiﬁers. Dried native caseins can also serve as food additive [4, 5]. Casein b and k can be separated from sodium caseinate. 3.1.1.2 Whey Proteins The main proteins are a-Lactalbumin (a-La, 14 kDa) and b-Lactoglobulin (b-Lg, 36 kDa in dimer form), which represent 70% of total whey proteins. a-La has several pharmaceutical applications and is added to infant milk while b-Lg can be used for emulsiﬁcation, foaming and gelling [6, 7] and can replace egg albumin in food products. It is also used as an additive in energetic drinks or in meat and ﬁsh basedproducts. Bovine serum albumin (BSA, 66 kDa) can be used for foaming and gelling in human food [8]. Lactoferrin (86 kDa) is used in cosmetics for skin protection and as anti-bacterial in meat preservative and in parenteral feeding [4, 7].

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Table 3.1 Average composition of cow milk: concentration and size distribution.

Concentration in whole milk (g/L) Water Fat globules Casein (in micelles) Serum proteins a-Lactalbumin b-Lactoglobulin BSA Proteose-peptone Immunoglobulins Lactoferrin Transferrin Others Lactose Mineral substances Organic acids Other

87.1 4.0 2.6 0.7 0.12 0.32 0.04 0.08 0.08 0.01 0.01 0.04 4.6 0.7 0.17 0.15

Size range and average (at weight average) 0.1–0.15 mm, average 3.4 mm 20–300 nm, average 110 nm 3–6 nm 14 kDa 18 kDa 66 kDa 4–40 kDa 150–900 kDa 86 kDa 76 kDa 0.35 kDa

3.2 Applications of Membrane Cross-Flow Filtration to Milk Processing 3.2.1 Milk Microfiltration

The main applications of MF to milk include bacteria and spore removal (cold pasteurization) and production of casein concentrates for milk standardization or cheese production with addition of cream. Milk is ﬁltered after its fat has been removed in order to avoid unnecessary membrane fouling. 3.2.1.1 Bacteria and Spore Removal This process does not heat denaturate whey proteins and provides longer preservation than pasteurization. However, it is necessary to transmit through the membrane all proteins, which is difﬁcult, due to the large micelle size and internal membrane fouling. A commercial process, Bactocatch, has been proposed by Alfa Laval (France), which consists [9] in combining large milk velocities (6–8 m s1) with a low uniform transmembrane pressure (TMP) in a ceramic tubular membrane with 1.4 mm pores. The uniform TMP is obtained by a cocurrent permeate recirculation with a pump to produce the same pressure gradient on both sides of the membrane and this process is known as UTP (or UTMP) mode [3]. Later, Isoﬂux tubular ceramic membranes with a continuous reduction in membrane thickness to reduce ﬁltration resistance along the membrane at the same rate as TMP have been proposed by TAMI Co (Nyons, France) [3]. SCT (now Exekia, Bazet, France) introduced Membralox ceramic

3.2 Applications of Membrane Cross-Flow Filtration to Milk Processing

membranes with a porosity gradient (GP) to achieve uniform ﬂux [3]. These two types of membranes do not require permeate recirculation and are therefore more economical in energy. Saboya and Maubois [10] reported a decimal log bacterial reduction of more than 3.5 with the Bactocatch system. 3.2.1.2 Casein Micelles Separation from Whey Proteins Unlike the case of bacterial removal, casein should be rejected by the membrane and pore sizes are smaller, from 0.2 to 0.05 mm, but the same type of systems with uniform TMP (UTP) or uniform ﬂux along the membrane at high ﬂuid velocity have been used for this application. Dauﬁn et al. [11] have used SCT membranes with 0.1-mm pores in the UTP mode with cocurrent permeate recirculation to separate caseins from whey proteins and obtained a whey-protein transmission of 70–80%. Gesan-Guiziou et al. [12], using a similar ceramic membrane (Kerasep 0.1 mm, TechSep Miribel, France) and the same ﬁltration bench in UTP mode reported ﬂuxes at 50 C of about 80 L h1 m2 with 50–80% a-La transmission, but permeate turbidity was relatively high (100–200 NTU), corresponding to about 2% casein transmission. Pouliot et al. [13] obtained permeate ﬂuxes of 90 L h1 m2 at ﬂuid velocity of 6.9 m s1 and a TMP of 190 kPa with a 0.22 mm pores Ceraﬂo ceramic membrane at a volume-reduction ratio (VRR) of about 1.5. Vadi and Rivzi [14] compared UTP and non UTP modes with a 0.2-mm pore ceramic Membralox multichannel membrane (Exekia, France). They obtained, in UTP mode, a ﬂux of 70 L h1 m2 at a VRR of 4, a TMP of 193 kPa, and a ﬂuid velocity of 7.2 m s1. They found that the non-UTP mode gave higher ﬂux up to a VRR of 4, while the UTP mode performed better at higher VRR. They also observed that the cake formed during MF in non-UTP mode was more difﬁcult to erode than the cake produced under UTP conditions. Le Berre and Dauﬁn [15] obtained a 99.5% casein retention at a ﬂux of 100 L h1 m2 with a 0.1-mm pore ceramic membrane and a whey-protein transmission between 70 and 90%. Samuelson et al. [16] used a 0.14-mm pore ceramic tubular membrane (Orelis, France) for casein concentration from skim milk, while minimizing whey-protein rejection by using cross-ﬂow velocities up to 8 m s1. They reported a maximum ﬂux of 145 L h1 m2 at a speed of 8 m s1 and 55 C, which fell to 80 L h1 m2 at 4 m s1. Whey-protein transmission was 88%, at 8 m s1 and 74% at 6 m s1, but casein rejection was low at 90%. A recent investigation of casein concentration by MF using polymeric membranes was made by Lawrence et al. [17] who used 0.3- and 0.5-mm pore PVDF (polyvinyliden ﬂuoride) membranes, both in a ﬂat-sheet laboratory module and in a spiral wound industrial pilot in non-UTP mode. They observed a casein rejection that increased from 96% at a TMP of 50 kPa to 98% at 150 kPa and 100% at 258 kPa. b-Lg transmission decreased from 22% at 50 kPa to 8% at 150 kPa and 1% at 258 kPa. In the ﬂat-sheet module at 50 C and a velocity of 0.44 m s1, the permeate ﬂux decayed from 60 L h1 m2 to 52 L h1 m2 over a period of 2 h. In the spiral module at the same velocity and 40 C, the ﬂux remained steady with time, at near 32 L h1 m2. Nelson et al. [18] developed a multistage MF process to remove a high percentage of whey proteins from skim milk while producing a low concentration factor retentate

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from microﬁltration. The microﬁltration retentate was blended with cream to standardize milk for traditional Cheddar cheese making. The MF permeate was ultraﬁltrated and the permeate obtained from this ultraﬁltration was diaﬁltered in order to remove whey proteins from skim milk before cheese making. The total process had 3 stages: the ﬁrst consisting in a MF of skim milk up to a VRR of 3, the second one was a ﬁrst diaﬁltration (DF) of permeate from ultraﬁltration and the last one was a second diaﬁltration. They used a UTP pilot (Tetra Alcross M7, Tetra Pack, Denmark) equipped with 0.1-mm pore ceramic membranes (Membralox). The TMP was maintained between 23–28 kPa. MF ﬂux was 30 L h1 m2. They removed about 95% of whey proteins. Zulewska et al. [19] microﬁltered pasteurized skim milk using several systems. The ﬁrst was a UTP pilot-scale with a ceramic 0.1 mm (Membralox, Pall Corp., East Hills, NY). The second was a 0.1-mm alumina membrane with graded porosity (GP, Membralox, Pall Corp.), and the third a polyvinylidene ﬂuoride (PVDF) spiral-wound (SW) module with 0.3-mm pores (Parker-Hanniﬁn, Tell City, Ind., USA) membranes. They found differences in ﬂux among ceramic UTP, ceramic GP, and polymeric SW microﬁltration membranes (54.08, 71.79, and 16.21 kg m2 per hour, respectively) when processing skim milk at 50 C in concentration tests until a concentration factor of 3 was obtained. These differences in ﬂux among the membranes would inﬂuence the amount of membrane surface area required to process a given volume of milk in a given time. The protein contents of microﬁltration permeates from UTP and GP membranes were higher than from SW membranes (0.57, 0.56, and 0.38%, respectively). Casein transmission in permeate was highest for the GP membrane and minimum in UTP module. The efﬁciency of removal of serum proteins was 64.40% in UTP mode, 61.0% and 38.6% respectively for GP and SW membranes. The SW polymeric membranes had a much higher rejection of serum proteins than the ceramic membranes. These data will be later compared with those obtained using dynamic microﬁltration. 3.2.2 Milk Ultrafiltration (UF)

Ultraﬁltration is used extensively in the dairy industry for concentrating proteins in cheese production by membrane [20, 21] and for the recovery of soluble proteins from whey [22]. A recently emerging application is the fractionation of whey proteins, mostly a-La and b-Lg [23, 24] for increasing their concentration in cheese or as food additives. This fractionation was previously achieved by chromatography, which gave a high purity, but a low output. 3.2.2.1 Total Proteins Concentration In order to retain, at least partially, a-La, the smallest whey protein, membranes must have a cut-off between 5 and 20 kDa. Clarke and Heath [24] have ultraﬁltered skim milk using 5 kDa polysulfone spiral-wound modules. Their permeate ﬂux was 14 L h1 m2 at 225 kPa and a cross-ﬂow velocity of 0.3 m s1. Labbe et al. [22]

3.2 Applications of Membrane Cross-Flow Filtration to Milk Processing

recovered and concentrated soluble proteins from whey by UF with a 20-kDa Carbosep membrane (zirconium oxide on carbon support, Techsep, Miribel, France). Permeate ﬂuxes were higher than for skim milk, but decayed during the ﬁrst hour of ﬁltration, due to protein–ZrO2 interactions. Yan et al. [25] ultraﬁltrated whole milk using tubular membranes (HBJ 180, Abcor Inc, USA). They obtained a maximum ﬂux of 42 L h1 m2 at 100 kPa, 49 C and a ﬂuid velocity of 3.13 m s1. The permeate ﬂux decayed linearly with VRR from 29 L h1 m2 at VRR ¼ 1 to 13 L h1 m2 at a VRR of 2.8. 3.2.2.2 Whey-Protein Fractionation Due to the difﬁculty of separating proteins with similar size such as a-La and b-Lg, most tests were not done on milk, but on binary protein mixtures or on protein concentrates. Cheang and Zydney [26] studied the separation of a-La and b-Lg from a binary mixture of these two pure proteins in a NaCl solution preﬁltered at 0.2 mm, using diaﬁltration (DF). This DF was performed with a small Amicon stirred-cell equipped with a 30-kDa cellulose membrane, at two pH of 5.5 and 7.2. With the 30-kDa membrane, a-La transmission was 26% at a permeate ﬂux of 12 L h1 m2 against only 0.5% for b-Lg. These transmissions increased with increasing ionic strength to reach 60% for a-La at a strength of 150 mM at pH ¼ 5.5, and 40% at pH ¼ 7.2. b-Lg transmissions were maximum at pH ¼ 7.2. Selectivity (ratio of a-La to b-Lg transmissions) reached a maximum of 58 at a pH of 5.5 and an ionic strength of 50 mM, but it decreased to 35 when permeate ﬂux was doubled. With a 50-kDa PES (Polyether sulfone) membrane, the maximum selectivity dropped to 10.5 at pH ¼ 5.5 and an ionic strength of 150 mM, due to the larger zeta potential of this membrane. The authors concluded that it was possible to separate a-La and b-Lg proteins with a high selectivity and a high yield rate, by optimal choices of pH, ionic strength and membrane cut-off. In a subsequent paper [27], the same authors obtained puriﬁed a-La and b-Lg fractions from whey protein isolate with a two-stage process. The ﬁrst step was a diaﬁltration at 100 kDa to separate a-La and b-Lg in permeate from BSA in retentate. The second step was an ultraﬁltration of permeate at 30 kDa followed by a DF in order to separate b-Lg in retentate from a-La in permeate. After 10 diavolumes, 75% of a-La was recovered in permeate. The ﬁnal selectivity was 21 at the end of second DF. They compared this process with a second one in which the ﬁrst DF was made at 30 kDa to collect a-La in permeate, while retentate was diaﬁltered at 100 kDa to collect b-Lg in permeate. This second process gave a higher a-La concentration than for the ﬁrst process, but a smaller yield, 85% instead of 95%. To produce puriﬁed a-La from acid casein whey, Muller et al. [28] proposed a prepuriﬁcation step by UF with a limited transmission of b-Lg. Membranes tested were a 150-kDa Carbosep M1 and ceramic ones (TAMI) of 150, 200 and 300 kDa. With the M1 membrane, a-La transmission decayed from 80% at 0.5 bar and a ﬂux of 30 L h1 m2 to 58% at 3 bar when permeate ﬂux rose to 80 L h1 m2. Transmissions were lower for the 300-kDa TAMI membrane and decayed with VRR from 35% at VRR ¼ 1.5 to 25% at VRR ¼ 4. They obtained a a-La yield in permeate of 53% and a purity (ratio of individual to total protein concentration) of 0.44 for a VRR of 9 with a

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permeate ﬂux of 30 L h1 m2. b-Lg transmission was 6% at a VRR of 3.5, and dropped to 4% at VRR ¼ 8. Their conclusion was that variations of physicochemical and hydrodynamic conditions could induce large differences in protein transmission. Almecija et al. [29] investigated the effect of pH (from 3 to 10) on the fractionation of whey proteins by diaﬁltration using a 300-kDa tubular ceramic membrane. a-La and b-Lg were collected in permeate while the retentate was enriched in BSA, immunoglobulins (Ig) and lactoferrins. Lowest permeate ﬂuxes were obtained at pH 4 and 5, the isoelectric point of a-La and b-Lg, due to increased fouling by aggregates of uncharged protein molecules, while the highest were obtained at pH 9 and 10, since membrane protein repulsion decreases aggregation and fouling. The largest yields of a-La in permeate (58%) were obtained at pH 7–9, and the lowest (4%) at pH 4. For b-Lg, the permeate yields followed the same trend, but were lower, 33% at pH 8 and 9 and 2% at pH of 4 and 5. Bramaud et al. [30] presented a process based on selective precipitation of a-La by heat treatment at 55 C for 30 min at pH of 3.9 followed by a centrifugation for separating in the soluble phase lactose and b-Lg from a precipitate containing BSA, Ig and a-La. In the second step, lactose was separated from b-Lg by diaﬁltration, at 0.5 mm while the precipitate was resolubilized with addition of CaCl2 to obtain a ﬁnal yield of 57% for a-La. Lucas et al. [31] obtained a maximum transmission of 37% for a-La and 10% of b-Lg, corresponding to a selectivity of about 3 using a 50-kDa Carbosep membrane. 3.2.3 Applications of Milk Nanofiltration (NF) and Reverse Osmosis (RO) 3.2.3.1 Treatment of Cheese Whey and Fabrication of Yogurts Cheese whey is generated by the traditional cheese fabrication consisting in coagulation of cream and casein. Each kilogram of cheese produces 5–10 kg of whey that contains about 6 g L1 of serum proteins, 48 g L1 of lactose and 613 g L1 of minerals. It is preferable to treat it as it constitutes a high COD (Chemical Oxygen Demand) efﬂuent and the proteins and lactose it contains can be recovered in the food and animal feed industry, after demineralization by electrodialysis or ion exchange. In order to save transportation costs, whey can be concentrated by RO or by evaporation before a two-stage treatment using UF to concentrate proteins in the ﬁrst retentate followed by NF to recover lactose in second retentate. Alternatively, a single NF step permits to concentrate serum proteins to 22% at VRR ¼ 4.5, while reducing the amount of minerals by 25–50% [32]. These serum proteins can be spraydried and used in various food applications under the names of whey protein concentrate (when containing 35–80% of proteins) or whey proteins isolates (with 80–95% of proteins) [3]. Nanoﬁltration has been used as an alternative to vacuum evaporation for concentrating milk in fabrication of yogurts, as it requires less energy. It is also used for selective demineralization of yogurts, for instance to lower sodium concentration or enrich them in magnesium or iron [20]. It is then possible to make low-fat yogurts

3.3 Dynamic Filtration

with better organoleptic properties than classical ones. But the main application of NFand RO seems to be the treatment of dairy process waters and efﬂuents, in order to recover milk proteins and lactose, while obtaining a depolluted permeate that can be recycled as water for rinsing or cooling if its ionic and lactose content has been sufﬁciently lowered. 3.2.3.2 Treatment of Dairy Effluents Dairy industry process waters resulting from starting, stopping or rinsing phases in the cheese-making process constitute a major source of milk protein loss as well as of pollution [33]. The chemical oxygen demand (COD) content of these efﬂuents, mainly due to the presence of lactose [34], is high, ranging generally from 500 to 6000 mgO2 L1. Most of the earlier work on this process has been done using NF or RO spiral-wound modules [34–36] because of their availability and relatively low cost. Balannec et al. [34], using milk diluted three times with an initial COD of 36 000 mgO2 L1 as an efﬂuent model with a spiral-wound module equipped with an Osmonics Desal 5 DL membrane of 150–300 Da cut-off. They obtained permeate ﬂuxes ranging from 24 L h1 m2 at initial concentration, a temperature of 25 C and a transmembrane pressure of 1900 kPa to 12 L h1 m2 at a volume-reduction ratio (VRR) of 5. The corresponding permeate COD rose from 125 mgO2 L1 at VRR ¼ 1 to 400 at VRR ¼ 5, remaining above the allowed French rejection limit of 125 mgO2 L1. Better COD removal was achieved when these authors used a Koch TFC HR reverse osmosis membrane that yielded a permeate COD of only 60 mgO2 L1 at VRR ¼ 5, but the corresponding permeate ﬂux fell from 18 L h1 m2 at VRR ¼ 1, to 7 at VRR ¼ 5. These permeate ﬂuxes were low because spiral-wound modules have a small hydraulic diameter (0.5 mm), and the high viscosity of concentrated milk prevented reaching high VRR. Vourch et al. [36] treated selected waste waters collected form dairy plants with a RO Koch TCR spiral-wound module in order to obtain recyclable water. Their permeate ﬂux decayed from 30 L h1 m2 at VRR ¼ 1 to 9 at VRR ¼ 5. They concluded that a RO þ RO cascade permitted to obtain a recovery 90–95% of water recyclable as boiler feed with a highly charged efﬂuent, against a single RO step for a low charged one. The total organic carbon in puriﬁed water was lower than 7 mg L1, against an initial value of 1000, while the conductivity was < 50 mS cm1.

3.3 Dynamic Filtration 3.3.1 Principle and Advantages of Dynamic (Shear-Enhanced) Filtration

We have seen in previous sections that in milk MF it was important to increase membrane shear rate by using high ﬂuid velocities while keeping TMP low and uniform, in order to transmit proteins through the membrane. This could only be achieved with permeate recirculation or specially designed membranes and the

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energy necessary to drive recirculation pumps was high. In whey protein fractionation by UF, the TMP had to be limited to retain sufﬁcient transmissions and permeate ﬂuxes were often low, from 25 to 30 L h1 m2. A RO stage was necessary to achieve sufﬁcient COD reduction in treatment of dairy process waters, leading to low ﬂux and high cost. Dynamic or shear-enhanced ﬁltration consists in creating the shear rate at the membrane by a disk rotating near a ﬁxed circular membrane or by rotating circular membranes around its axis or by vibrating the membrane either longitudinally or torsionally around a perpendicular axis [37]. This mode of ﬁltration can generate very high shear rates at the membrane that not only increase substantially the permeate ﬂux, but have a favorable effect on membrane selectivity. Microsolute transmission is increased in dynamic microﬁltration, which reduces cake formation by combining a high shear rate with a low TMP. In addition, high shear rates reduce concentration polarization and the concentration of rejected solutes at the membrane. Thus, concentration gradient and diffusive solute transfer through the membrane are decreased, which increases solute rejection rates in NF and RO, when mass transfer through the membrane is mainly diffusive. At the same time, permeate ﬂuxes keep increasing until high pressures, as the pressure-limited regime is extended by the reduction of concentration polarization and very high ﬂuxes can be obtained at high TMP. The inlet ﬂow rate into the module needs to be only slightly larger than the ﬁltration ﬂow rate, reducing pumping energy. The drawbacks of dynamic ﬁltration are its complexity and limited membrane area for some systems, such as multicompartment rotating-disk systems, which raise the equipment cost. But, the recent availability of large-diameter ceramic disk membranes permits the construction of immersed rotating membranes of 80 m2 area or more in a single housing, which are easier and less costly to build than multicompartment systems. 3.3.2 Industrial Dynamic Filtration Systems

The ﬁrst commercialized dynamic ﬁltration systems were of Couette ﬂow type with cylindrical membranes rotating inside a concentric cylindrical housing, such as the Biodruckﬁlter (Sulzer AG, Winterthur, Switzerland) and the Benchmark Rotary Bioﬁltration (Membrex, Garﬁeld, NJ, USA) [38]. This concept takes advantage of Taylor vortices created at large speed in the annular space between membranes and housing that increase the shear rate, but the maximum membrane area of commercial systems is about 2 m2. The Dyno system, manufactured by Bokela GmbH (Karlsruhe, Germany), consists in several disks rotating on the same shaft between ﬁxed circular membranes for a total membrane area up to 8 m2. Its maximum pressure is 600 kPa (Figure 3.1). It is available with polymeric or ceramic (metallic) membranes. The Optiﬁlter CR (Metso Paper Raisio, Finland) features blades rotating between stationary ﬂat circular membranes with a tip azimuthal speed of 10 to 15 m s1. Its total membrane area can exceed 140 m2 with a 132-kW motor [39]. They are used

3.3 Dynamic Filtration

Figure 3.1 Dyno rotating-disk module (Bokela, Germany).

by more than 30 plants, mostly for treatment of pulp and paper efﬂuents or pigment recovery. The recent availability of ceramic membrane disks, especially in Germany, has spurred the commercialization of multishaft systems with overlapping rotating membranes. For instance, the MSD (Multi Shaft Disk) system (Westfalia Separator, Aalen, Germany) features 31.2 cm diameter ceramic membranes mounted on 8 parallel shafts arranged as shown in Figure 3.2. All disks rotate at the same speed and are enclosed in a cylindrical housing. Other systems, the Rotostream (Canzler, Dueren, Germany) [40] and the Hitachi (Tokyo, Japan) available up to, respectively, 150 and 100 m2 membrane area have their parallel axes in the same plane. The Novoﬂow Company, (Oberndorf, Germany) manufactures single-shaft rotating MF and UF ceramic membranes systems, the SSDF (Single Shaft Disk) using 312-mm ceramic disks for a membrane area of 15 m2 per module. The company reported a low energy consumption of 2.5 kW for a 15-m2 module, corresponding to 0.64 D per m3 of permeate and a total operating cost of 7.4 D /m3. The SSDF is also available with composite MF-UF-NF membranes of 55 cm diameter with 25 m2 of membrane per module. Krauss-Maffei Process Technology (KMPT AG, Germany, www.kmpt.com), has developed a dynamic ﬁltration module, similar to the MSD, but which can be equipped with rotating ceramic or polymer membrane disks. The module is in stainless steel and has a membrane area of up to 16.4 m2. Membrane pore sizes range from 7 nm to 2 mm. A vibratory membrane system (VSEP, New Logic Emeryville, Ca, USA), consists of a stack of circular organic membranes (Figure 3.3), mounted on a vertical torsion shaft spun in azimuthal oscillations by a vibrating base, at its resonant frequency of about 60 Hz. The shear rate at the membrane is produced by the inertia of the retentate that moves at 180 out of phase with the membrane and varies sinusoidally with time. The use of resonance minimizes the power necessary to produce the vibrations, which is only 9 kW, even for large units of 150 m2 membrane area (Figure 3.4) The key parameter governing performance is the maximum azimuthal displacement of the

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Figure 3.2 Industrial MSD module with 8 parallel shafts and 31-cm ceramic disks. Courtesy of Westfalia Separator.

Figure 3.3 Schematic of circulation in VSEP membrane stack (Courtesy of New Logic Research).

3.3 Dynamic Filtration

Figure 3.4 Industrial VSEP vibrating modules (Courtesy of New Logic Research).

membrane rim, which has been measured as a function of frequency in [41] and is limited to about 3 cm. The VSEP has been used for the ﬁrst time in Europe in 2007 to treat anaerobically digested pig manure. The system was installed and commissioned in Belgium at a major pig farm where it will be used for the biomethanation of raw manure, a comprehensive process developed by the Belgian ﬁrm where methane is recovered and converted into electrical energy. Zouboulis and Petala [42] studied the performance of VSEP for the treatment of raw stabilized leachate produced during landﬁll of municipal wastes. Four different membrane types were examined for the treatment of leachates, that is, one for microﬁltration (0.1 mm), two for ultraﬁltration (100 and 10 kDa) and one for nanoﬁltration (50% rejection of NaCl). The removal of organic matter in terms of COD value exceeded 60% for all cases. The PallSep (Pall Corp, USA) is Palls version of the VSEP intended for biotechnological and food applications and is available with up to 32 m2 of membrane area. Postlethwaite et al. [43] investigated this system for protein recovery from a model biological feed stream containing 200–500 g L1Saccharomyces cerevisiae and 0.75 g L1 bovine serum albumin (BSA). They reported that the ﬂux and transmissions at a biomass concentration of 500 g L1, were 45 L h1 m2 and 67%, respectively, and could be maintained over extended periods. 3.3.3 Application of Dynamic Filtration to Skim-Milk Processing 3.3.3.1 Casein Separation from Whey Proteins by MF One of the ﬁrst applications of dynamic ﬁltration to this task in UHT milk has been made with a VSEP pilot [44] equipped with a 500-cm2, 0.1-mm pore Teﬂon membrane using UHT milk. The permeate ﬂux at 45 C and maximum vibration frequency (60.75 Hz) reached a plateau of 95 L h1 m2 at 100 kPa. This ﬂux decayed with time to 50 L h1 m2, which corresponded to the critical ﬂux for stable operation. Permeate turbidity decayed with time from 52 to 15 NTU, indicating very good casein micelle

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Figure 3.5 Variation of permeate flux and frequency with VRR in MF of powder milk (from Ref. [45] with permission).

rejection. Similar tests, performed with the same VSEP pilot and membrane, but using powder skim milk with same protein composition as pasteurized milk have been reported in [45]. In concentration tests at 55 kPa, the ﬂux decayed from 50 to 33 L h1 m2 at VRR ¼ 2 (Figure 3.5). The faster initial rate of decay is due to the lower frequency of 60.2 Hz. When the 0.1-mm membrane was replaced by a 150-kDa PES one, the permeate ﬂux decayed slowly with increasing concentration from 40 L h1 m2 to 35 at VRR ¼ 2, (Figure 3.6) while permeate turbidity dropped from 160 NTU to about 30 indicating good micelle rejection (Figure 3.6). When TMP was varied over a cycle, the permeate ﬂux was reversible, but a-La transmission, which was between 70 and 80%, and b-Lg (30–35%) decayed with time

Figure 3.6 Variation of permeate flux and turbidity with VRR in ultrafiltration of powder milk (from Ref. [45] with permission).

3.3 Dynamic Filtration

Figure 3.7 Variation of UF permeate flux and whey-protein transmission with TMP during a pressure-variation cycle (from Ref. [45] with permission).

(Figure 3.7) due to internal fouling. Espina et al. [46] microﬁltered skim UHT milk using a MSD pilot with six 9-cm diameter rotating ceramic membranes with 0.2-mm pores. Permeate ﬂuxes reached a maximum of 120 L h1 m2 at a rotation speed of 1930 rpm, a TMP of 100 kPa, and 40 C (Figure 3.8). Permeate turbidity was less than 20 NTU, indicating excellent casein micelles rejection. In concentration tests, the permeate ﬂux decayed logarithmically with VRR (Figure 3.9) according to the thinﬁlm theory of Blatt et al. [47]. The reduction from 1930 to 1044 rpm has a large effect on permeate ﬂux. Corresponding a-La and b-Lg transmissions, are shown in Figures 3.10 and 3.11 140 MSD; VRR = 1; T= 40°C; Q = 3 L/min; 0.2 µm ceramic.

120

J (L h–1 m–2)

100 80 60 40 1931 rpm 1492 rpm 1044 rpm 745 rpm

20 0 0

50

100

150

200

250

300

350

TMP (kPa) Figure 3.8 Variation of stabilized permeate flux versus TMP with the MSD at different rotation speeds for tests of Figure 3.4 and one at 1930 rpm (from Ref. [46], with permission).

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1

0.9 0.8

Tr α-La

0.7 0.6 0.5 0.4 Test 1; 1044 rpm Test 2; 1930 rpm Test 3; 1930 rpm

0.3 0.2 0.1

MSD. Ceramic membranes 0.2 µm. Q = 3 L/min. T = 40°C

0

1

1,5

2

2,5

3

3,5

4

4,5

VRR Figure 3.9 Variation of permeate flux with VRR (semi-log) in MF of skim milk with the MSD module for tests 1 to 3 (from ref. [51], with permission).

respectively. At a speed of 1930 rpm, these transmissions remain between 80 and 90% after about 15 min of ﬁltration until the maximum VRR. At 1044 rpm, these transmissions reach a maximum at VRR ¼ 1.3 and decrease at higher VRR to 50% for a-La and 40% for b-Lg. The same group also tested a prototype rotating-disk module, designed at the University of Technology of Compiegne (UTC), consisting in a metal disk equipped with radial vanes rotating at high speed near a ﬁxed 0.15-mm pore PVDF circular membrane. This module yielded higher ﬂuxes, up to 200 L h1 m2 at a speed of 2000 rpm and 200 kPa (Figure 3.12) since the membrane shear rate was higher than in the MSD due to the larger membrane radius (15 cm instead of 9). Permeate turbidity was also very low at 10 NTU, indicating casein rejection higher 120 Test 1; 1044 rpm Test 2; 1930 rpm Test 3; 1930 rpm

J (L h–1 m–2)

100

JTest 1 = –23.91Ln(VRR) + 53.88 R 2 = 0.994 JTest 2 = –47.04Ln(VRR) + 103.23 X

80

R 2 = 0.988 JTest 3 = –48.95Ln(VRR) + 106.68 R 2 = 0.994

60

40

20 MSD. Ceramic membranes 0.2 µm. Q = 3 L/min. T = 40°C

0 1

VRR

10

Figure 3.10 a-La transmission with VRR for MF tests of Figure 3.9. (from Ref. [51], with permission).

3.3 Dynamic Filtration 1 0.9 0.8

Tr β-Lg

0.7 0.6 0.5 0.4 0.3 Test 1; 1044 rpm Test 2; 1930 rpm Test 3; 1930 rpm

0.2 0.1

MSD. Ceramic membranes 0.2 µm. Q = 3 L/min. T = 40°C

0 1

1,5

2

2,5

3

3,5

4

4,5

VRR Figure 3.11 b-Lg transmission with VRR for MF tests of Figure 3.10 (from Ref. [51], with permission).

than 99.5%. However, a-La and b-Lg transmissions were low, respectively 30–35% for a-La and 8% for b-Lg (Figure 3.13) due in part to the lower membrane cut-off. These data conﬁrmed the high potential of rotating disks and rotating membrane systems that performed better than the VSEP for this application. As seen in Section 3.3, permeate ﬂuxes with tubular ceramic membranes in UTP mode were generally between 70 and 90 L h1 m2 at 50 C with tangential velocities of about 7 m s1 and casein micelles rejection was generally not as high as with the MSD. a-La and b-Lg

270 Rotating disk with 6 mm vanes; 2000 rpm; T = 45°C; Q = 3 L/min; 0.15 µm PVDF

240

J (L m–2 h–1)

210 180 150 120 90 New membrane 1st reuse 2nd reuse 3rd reuse

60 30 0 0

50

100

150

200

250

300

350

TMP (kPa) Figure 3.12 Variation of stabilized permeate flux versus TMP using the rotating-disk module with vanes at 2000 rpm and with a new 0.15-mm PVDF membrane and after several reuses (from Ref. [46], with permission).

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j 3 Milk and Dairy Effluents Processing: Comparison of Cross-Flow and Dynamic Filtrations

60

0.5

Rotating disk with 6 mm vanes; 2000 rpm; T = 45°C; Q = 3 L/min; 0.15 µm PVDF

0.3 0.25

0.4

0.15 0.2

Tr β-Lg

Tr α-La

0.2 0.3

0.1 0.1

0.05

0 0

50

100

150

200

250

300

0 350

TMP (kPa) Figure 3.13 Variation of a-La and b-Lg transmissions with TMP for the tests of Figure 3.12 (from Ref. [46], with permission).

transmissions obtained with the MSD pilot compared favorably with those reported in UTP mode [13]. 3.3.3.2 Dynamic Ultrafiltration of Skim Milk Jaffrin et al. [48] compared the performance of rotating disk and VSEP modules equipped with the same PES 50-kDa membrane. Two types of disks were tested, a ﬂat (or smooth) disk and a disk equipped with eight 6-mm high radial vanes. Due to reduced concentration polarization by high shear rates, the permeate ﬂux kept rising with increasing TMP for the disk with vanes that produces a maximum shear rate at disk periphery of 2.8 105 s1, until at least 600 kPa, reaching 200 L h1 m2 (Figure 3.14). With the same disk rotating at 1000 rpm, the maximum membrane shear rate fell to 8.2 104 s1, which was about the same as for a smooth disk rotating at 2000 rpm. Permeate ﬂuxes for these two cases were almost the same, reaching 115 L h1 m2 at 400 kPa. The VSEP, which had a slightly higher shear rate of 1.15 105 s1, reached the same ﬂux, but at a higher TMP of 850 kPa. The same comparison, but made during concentration tests is shown in Figure 3.15. The highest permeate ﬂuxes were obtained with a disk equipped with vanes rotating at 2000 rpm, which is logical, since it corresponds to the maximum shear rate. The permeate ﬂux decayed slowly, from 130 to 120 L h1 m2 until VRR ¼ 3, as it is pressure limited. Then it dropped at a faster rate as Ln (VRR1) as the ﬂux became mass transfer limited at higher VRR, since the increase in viscosity lowered the shear rate. When a smooth disk was used at the same speed, the ﬂux was lower and masstransfer limited as the membrane shear rate was one third of the previous case. The VSEP permeate ﬂux was slightly lower than for the smooth disk for VRR < 3.5, even though TMP was higher, 400 kPa instead of 300. For VRR > 3.5, however, the VSEP ﬂux exceeded that of the rotating disk, since the VSEP shear rate decreases less at high concentration than with the rotating disk.

3.3 Dynamic Filtration 250 UHT Skim milk, 50 kDa PES membrane, VRR=1, T=45 °C

225 200

γm = 2.84×105 s–1 Disk with 6 mm vanes, N = 2000 rpm Disk with 6mm vanes, N = 1000 rpm Smooth disk, N= 2000 rpm Smooth disk, N= 1000 rpm Vsep, F=60.75 Hz

J (l.h–1.m–2)

175 γm = 8.17×10 s

4 –1

150 125 100 75 50

γm = 1.15×105 s–1

γm = 9.12×104 s–1

γm = 2.62×104 s–1

25 0 0

100

200

300

400

500

600

700

800

900

1000

TMP (kPa) Figure 3.14 Variation of permeate flux in UF of skim milk versus TMP using the rotating-disk module with a smooth disk and a disk with vanes, and the VSEP (from ref. [48], with permission).

Ding et al. [49] ultraﬁltered UHT milk with the same PES 50 kDa membrane as in [48] using a rotating-disk module. They measured the net power (PN) consumed by friction on the disk as function of rotation speed (N) together with corresponding permeate ﬂow rates QF (Figure 3.16). Since in a small pilot, the power consumed by the shaft and internal parts of motor is disproportionably high, the power consumed by the motor with an empty module was subtracted from the power measured at the same speed during milk ﬁltration, in order to obtain power consumed by disk friction 50 kDa PES membrane, UHT Skim milk, T= 45 °C

140 120

J (L.h–1.m–2 )

100

J = –58.9 Ln(VRR) + 110.1 ; N=2000 tr/min, TMP = 300 kPa

J = –104.4 Ln (VRR) + 231.8 ; N=2000 rpm, TMP = 300 kPa

80 60 J = –40.6 Ln (VRR) + 87.7 ; F=60.75 Hz, TMP = 400 kPa

40 Disk with 6 mm vanes, small module smooth disk, small module Vsep

20 0 1

VRR

10

Figure 3.15 Variation of permeate flux in UF of skim milk versus VRR in semi-logarithmic scale using the rotating-disk module with two types of disks and the VSEP (from ref. [48], with permission).

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j 3 Milk and Dairy Effluents Processing: Comparison of Cross-Flow and Dynamic Filtrations

62

500

6

UHT Skim milk, 50 kDa PES, T=45 °C

450 5 PN

350

PN (W)

QF

300

disk with 6-mm vanes smooth disk disk with 6-mm vanes smooth disk

4

250

3

200

Q F (L.h-1)

400

2

150 100

1

50 0 0

500

1000

1500

2000

2500

0 3000

N (rpm) Figure 3.16 Variation of net power consumed by the disk PN and permeate flow rate QF for two types of disks with rotation speed (from ref. [49], with permission).

alone, which will be the dominant part in a large module. As expected, the power increased as N2 and was larger for a disk with vanes and the gap between the two disks widened at large speed. The speciﬁc power per m3 of permeate, plotted in Figure 3.17, which is given by the ratio PN/QF, increased with N and was higher for a disk equipped with vanes than for a smooth one, as the increment in permeate ﬂow rate with vanes was less than the power increase. But vanes increase the ﬂux and permit to lower membrane area. Thus, higher energy costs may be offset by a reduction in

160 UHT Skim milk, 50 kDa PES, T=45 °C

140

–3

PN/Q F (kWh.m )

120 100

Disk with vanes, s = 6 mm Smooth disk

80 60 40 20 0 0

500

1000

1500

2000

2500

3000

N (rpm) Figure 3.17 Variation of specific energy consumed by the disk per m3 of permeate using data of (from ref. [49], with permission).

3.3 Dynamic Filtration

equipment cost. Optimal conﬁguration and rotation speed may be determined from appropriate ﬁnancial and economic information. 3.3.3.3 Total Protein Concentration by UF for Cheese Manufacturing Akoum et al. [50] used a VSEP pilot equipped with a 10-kDa PES membrane permitting high protein rejection to concentrate caseins and whey proteins from a powder low-heat skim milk with the same composition as fresh milk. Permeate ﬂuxes obtained at 46 C and initial concentration are given in Figure 3.18 for various vibration frequencies. In order to reduce wear and maintenance, the VSEP is not used at its maximum frequency in normal industrial use, but with a membrane displacement amplitude of 2–2.5 cm at the rim, rather than at the maximum of 3 cm at resonance. This corresponds to frequencies of 60–60.2 Hz for this pilot. The permeate ﬂux kept increasing with TMP until 1500 kPa, even at 60 Hz where it reached 70 L h1 m2, while at lower frequencies the maximum was reached at 600 kPa or less. Variations of permeate ﬂuxes in concentration tests without permeate recycling are displayed in Figure 3.19 and decay linearly with increasing Ln(VRR). The maximum theoretical VRR, extrapolated to zero ﬂux, was about 17 for all frequencies, thus, higher than corresponding values obtained with cross-ﬂow ﬁltration, which are less than 10. A comparison of variation of permeate ﬂuxes versus milk dry mass in %, which is proportional to the concentration factor, is shown in Figure 3.20 for UHTand powder skim milks and for 10- and 50-kDa membranes. Data for powder and UHT milks are very close although, in UHT milk, whey proteins are partially denatured and the ﬂux dropped a little faster with increasing concentration for the 50-kDa membrane, due perhaps to larger internal fouling.

100 F (Hz) 60.75 60.4 60.2 60 59.75 59.5 59 58 57 55 0

90 80

J (L.h–1.m–2)

70 60 50 40 30

Powder milk, 10 kDa PES Membrane, T=46 °C

20 10 0 0

200

400

600

800

1000

1200

1400

1600

TMP (kPa) Figure 3.18 Variation of permeate flux in UF at 10 kDa of skim milk with TMP using a VSEP at various frequencies and VRR ¼ 1 (from ref. [50], with permission).

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j 3 Milk and Dairy Effluents Processing: Comparison of Cross-Flow and Dynamic Filtrations

64

120 Powder milk, T=45 °C, 10-kDa PES Membrane

–1

–2

J (L.h .m )

100 F (Hz)

80

60

2

60.75

J = –34.3 Ln(VRR) + 96.2 , R = 0.991

60.4

J = –27.0 Ln(VRR) + 77.4 , R = 0.997

60.2

J = –25.5 Ln(VRR) + 72.3 , R = 0.996

2

2

40 60

2

J = –23.2 Ln(VRR) + 67.0 , R = 0.991

20 17

0 1

10

20

100

VRR Figure 3.19 Variation of permeate flux in UF at 10 kDa of skim milk with VRR using a VSEP at various frequencies and a TMP of 1.5 MPa (from ref. [50], with permission).

3.3.3.4 a-La and b-Lg Protein Fractionation by UF Espina et al. [51] used a UTC rotating-disk module equipped with a 50-kDa PES membrane on skim UHT milk permeate obtained after MF at 0.2 mm with ceramic membranes to separate a-La in UF permeate from b-Lg in retentate. The UF permeate ﬂux obtained at 40 C, shown in Figure 3.21 was higher, at 200 L h1 m2 and VRR ¼ 4, than those reported in Section 3.2 with cross-ﬂow UF, which were less than 100 L h1 m2. a-La transmission, shown in Figure 3.22, rose from a minimum of 11% at VRR ¼ 1.3–24% at the maximum VRR of the test (3.1). b-Lg transmission

90 T=45 °C, F=60.75 Hz

80

J (L.h–1.m–2)

70 2

60

J = –65.6 Ln (DM) + 228.9 ; R = 0.9866 50-kDa PES Membrane

50 2

J = –52.4 Ln(DM) + 191.07; R = 0.9945 10-kDa PES Membrane

40 30

Powder Milk 50 kDa

20

UHT Milk 50 kDa

10

Powder Milk 10 kDa UHT Milk 10 kDa

0 1

10

20

30

40

100

Dry Mass, DM (%) Figure 3.20 Variation of permeate flux in UF at 10 and 50 kDa of UHT and powder skim milks with dry mass percentage using a VSEP (from ref. [50], with permission).

3.3 Dynamic Filtration 300

200

–1

–2

J (L h m )

250

150

100

50

Rotating-disk module, disk with vanes. 50-kDa PES membrane. 2000 rpm. Q i = 3 L/min. 40 °C

0 1

10

VRR

Figure 3.21 Variation of permeate flux with VRR for UF at 50 kDa of UHT milk MF permeate using the rotating-disk module at 2000 rpm (from ref. [51], with permission).

dropped to about 2–3% at VRR > 1.8, so that selectivity (Tra/Trb), also shown in Figure 3.22, rose to 8 at VRR ¼ 3.1. This selectivity was close to that of 10.5 obtained with a 50 kDa membrane, but on a binary protein mixture by Cheang and Zydney [26], after optimizing ionic force and pH. By contrast, Gesan-Guiziou et al. [52] obtained a transmission of only 9% for a-La and 6% for b-Lg during the ultraﬁltration of redissolved precipitate from Gouda whey protein concentrate with a 50-kDa Carbosep membrane, at VRR ¼ 10 and 50 C.

0.3

9

Rotating-disk module, disk with vanes. 50-kDa PES membrane. 2000 rpm. Q i = 3 L/min. 40 °C

8

Tr α-La; β-Lg

7 6

0.20

5 0.15 4 α -La a-La β -Lg b-Lg

0.10

3

S

2

Selectivity (Tr α-La/Tr β-Lg)

0.25

0.05 1 0

0 1

1,5

2

2,5

3

3,5

VRR Figure 3.22 a-La and b-Lg transmissions, and variation of selectivity (Tr a-La)/(Tr b-Lg) versus VRR for test of Figure 3.21 (from ref. [51], with permission).

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66

Bhattacharjee et al. [53] separated b-Lg from whey protein concentrate obtained from raw casein whey by centrifugation followed by a MF at 0.45 mm. Their dynamic ﬁltration module consisted in a circular polymer membrane of 76 mm diameter rotating inside a cylinder, near a disk stirrer rotating in the opposite direction at 500 rpm. They used a complex three-stage process, starting with a diaﬁltration at 5 kDa to remove lactose, minerals and salts. The retentate was then ultraﬁltered at 30 kDa, after addition of hydrochloric acid to lower the pH to 2.8 in order to obtain monomer b-Lg and a-La, while bovine serum albumin, lactoferrin and immunoglobulins were collected in retentate. When the membrane was at rest, the ﬂux decayed from 200 L h to 20 L h1 m2 after 20 min of ﬁltration. When the membrane speed was set to 300 and 600 rpm, the ﬂux stabilized to 100 and 115 L h1 m2, respectively. The ﬁnal separation between monomer b-Lg and a-La was obtained by ion-exchange membrane chromatography as the molecular weights of these two proteins were too close to be separated by UF. The separation factor between b-Lg and a-La increased with the pH of the loading buffer in ion–exchange chromatography to reach a maximum of 4.7 at pH ¼ 5.0. The ﬁnal purity of b-Lg, relative to total proteins, was 0.87. The lowering of pH to 2.8 permitted to increase the b-Lg/a-La ratio to 17.15 as compared to 9.64 when b-Lg remained in dimer form at pH ¼ 5.6. 3.3.4 Treatment of Dairy-Process Waters by Dynamic NF and RO

Akoum et al. [54] used a L101 VSEP pilot to treat white waters represented by one volume of skim UHTmilk diluted with two volumes of pure water. The initial COD of this diluted milk, mainly due to lactose, was 36 000 mgO2 L1, which corresponds to a highly charged efﬂuent. The VSEP was equipped with the same Desal 5DK and 5DL as spiral-wound modules used by Balannec et al. [34]. Variations of permeate ﬂux, COD and conductivity (proportional to ion concentration) obtained using the VSEP at 25 and 45 C are represented in Figure 3.23 as a function of TMP for a 5DL membrane and initial concentration. For comparison, the graph also indicates permeate ﬂux and COD provided by the spiral module equipped with the same 5 DL membrane at 25 C and 1.9 MPa. The spiral module ﬂux was 24 L h1 m2 or one third of the VSEP ﬂux at same TMP and temperature (72 L h1 m2). The spiral module COD was 128 mgO2 L1, ﬁve times higher than the VSEP COD (24 mgO2 L1) under the same conditions. It is interesting that the high shear rates of the VSEP, not only increase signiﬁcantly the permeate ﬂux as compared to cross-ﬂow ﬁltration, but decrease lactose and ions transmission, responsible for permeate COD in NF by reducing their concentration at the membrane due to lower concentration polarization. In concentration tests without permeate recycling (Figure 3.24), the VSEP retains its high performance with a permeate ﬂux which decayed linearly with increasing VRR to 30 L h1 m2 at VRR ¼ 5, 1.9 MPa and 25 C, against 11 L h1 m2 for the spiral-wound module under same conditions. Presumably, the ﬂux difference between the two modules would have been larger at higher TMP as the VSEP ﬂux kept increasing until TMP ¼ 4 MPa, while the spiral-module one leveled off at about 2 MPa. However,

3.3 Dynamic Filtration 1500

Flux at 25 °C COD at 25 °C Flux at 45 °C COD at 45 °C Spiral module flux Spiral module COD Conductivity 25 °C Conductivity 45 °C

200

5 DL membrane, F=60.75 Hz, VRR=1 1200

150

900

100

600

50

300

0

Conductivity (μS.cm–1)

J (L.h–1.m–2),COD (mgO2.L–1)

250

0 0

0,5

1

1,5

2

2,5

3

3,5

4

4,5

5

TMP (MPa) Figure 3.23 Variation of permeate flux, permeate COD and conductivity of diluted milk with TMP using a VSEP and a nanofiltration membrane and comparison at 25 C with a spiral-wound module with the same membrane and conditions (from ref. [54], with permission).

VSEP COD, which was half that of spiral module up to VRR ¼ 2, increased faster at high VRR and COD of both modules reached 350 mgO2 L1at VRR ¼ 5. A similar investigation, but using a rotating-disk module, was carried out with the same model efﬂuent (diluted milk) and a Desal 5 DK membrane at a temperature of 45 C and TMP of 4 MPa by Frappart et al. [55]. Variations of permeate ﬂux with VRR at rotation speeds of 1000 and 2000 rpm and two types of disks are presented in Figure 3.25. As expected, the highest permeate ﬂuxes were obtained with a disk equipped with 6-mm radial vanes and rotating at 2000 rpm, producing a maximum 100

1500 Flux Spiral module at 1.9 MPa, 5 DL and 25 °C Flux VSEP at 2 MPa and 60.75 Hz, Desal 5 DK and 25 °C COD Spiral module at 1.9 MPa COD Standard COD VSEP at 2 MPa and 60.75, 5 DK

J (L.h–1.m–2)

80

1350 1200

70

1050

60

900

50

750

40

600

30

450

20

300

10

150

0

COD (mgO2.L–1)

90

0 1

2

3

4

5

6

7

8

VRR Figure 3.24 Comparison of permeate flux and COD variations with VRR in NF. (from ref. [54], with permission)

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j 3 Milk and Dairy Effluents Processing: Comparison of Cross-Flow and Dynamic Filtrations

68

250 2000 rpm ; 6-mm vanes 1000 rpm ; 6-mm vanes 2000 rpm ; smooth 1000 rpm ; smooth 2000 rpm ; 6-mm vanes ; undiluted milk

VRR max = 12.7 VRR max = 13.6 VRR max = 10.5

TMP = 4000 kPa ; T = 45 °C

150

–1

–2

J (L h m )

200

VRR max = 14.8

100

50

0 1

10

100

VRR Figure 3.25 Variation of permeate flux with VRR at a TMP of 4 MPa using a rotating-disk module at 1000 and 2000 rpm for two types of disk and a 5DK NF membrane (from ref. [55], with permission).

shear rate at membrane periphery of 4.4 105 s1. This ﬂux decayed from 225 L h1 m2 at VRR ¼ 1 to 140 at VRR ¼ 5, while with the same disk rotating at 1000 rpm or with a smooth disk rotating at 2000 rpm, the ﬂux at VRR ¼ 5 dropped to about 90 L h1 m2 as respective shear rates were only 1.2 105 and 1.1 105 s1. Corresponding variations of permeate COD with VRR are represented in Figure 3.26. These COD are lowest at the highest shear rates, but they exceed the allowed limit (rejection standard) of 125 mgO2 L1 above VRR ¼ 4, so that a RO step may be 800 2000 rpm ; 6-mm vanes

Permeate COD (mgO2 L–1)

700

1000 rpm ; 6-mm vanes 2000 rpm ; smooth

600

1000 rpm ; smooth COD rejection standard

500

TMP = 4000 kPa; T = 45 °C

400 300 200 100 0 0

1

2

3

4

5

6

7

8

9

VRR Figure 3.26 Variations of permeate COD with VRR for the tests of Figure 3.25 (from ref. [55], with permission).

3.4 Conclusion

necessary at high VRR. However, for usual dairy efﬂuents with lower initial COD, the limit may be respected with a single NF step.

3.4 Conclusion

The use of membrane processes in the dairy industry has increased signiﬁcantly during the last 20 years. Bacteria removal by MF avoids serum protein denaturation and nutritional losses due to UHT or pasteurization treatment. Recently available Isoﬂux membranes with permeability gradient are replacing UTP processes with cocurrent permeate recirculation that required more energy. Membrane processes for protein fractionation are emerging as they can be extrapolated to large volumes and automatized production, unlike ion exchange, afﬁnity chromatography and selective precipitation. According to Brans et al. [3], their technical advantages should spur their industrial development. But milk is a complex ﬂuid that presents a challenge to membrane processes, as many of its components induce fouling, which requires use of large ﬂuid velocities and highly selective membranes. Thus, process conditions and fouling control methods must be further optimized. Dynamic ﬁltration, which has clearly proved its efﬁciency to reduce membrane fouling in MF and concentration polarization in UF, NF and RO, may play an important role, especially for extracting valuable milk components. Systems with rotating ceramic membranes and vibrating ones seem well suited for this application, but their costs are presently higher than those with tubular membranes, because of their small production. But their cost should decrease as sales increase. In addition, dynamic ﬁltration gives the choice between using high shear rates with large rotation speeds in order to increase permeate ﬂux, or to use moderate rotation speeds giving the same ﬂux as in cross-ﬂow ﬁltration, but with a lower energy per m3 of permeate. Thus energy savings may compensate the higher initial cost. Dynamic microﬁltration with rotating ceramic disks may be another alternative to cocurrent permeate recirculation. Concerning applications involving NF and RO, the industrial future of dynamic ﬁltration is more delicate to predict. This chapter has clearly shown the high performance of VSEP and rotating-disk modules, both equipped with polymer membranes, as no NF and RO ceramic disks seem to be yet available. As said earlier, polymeric-membrane modules for dynamic ﬁltration are more complex and costlier to build than ceramic membranes modules. In addition, large spiral-wound modules, which are built in large quantities for water desalination and water treatment are very inexpensive, about 10–15 D per m2 and are also very compact. So dynamic modules probably cannot presently compete with spiral-wound NF and RO modules in terms of cost per m3 of permeate, even if their membrane area is much smaller for the same output. The situation may be different for fractionation applications that are generally carried out with tubular ceramic membranes of much higher cost than spiral-wound modules, and for which a high selectivity and a low energy consumption are important.

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70

Acknowledgments

V. Espina has been supported by a scholarship No. E06D101610CL of the Alban European Union Program for Latin America. The authors thank Westfalia Separator for the loan of a MSD pilot, and Vivendi Waters for the loan of a VSEP pilot.

Abbreviations

DF MF NF RO TMP UF UHT UTC UTP VRR

diaﬁltration microﬁltration nanoﬁltration reverse osmosis transmembrane pressure ultraﬁltration ultrahigh temperature University of Compiegne uniform transmembrane pressure volume-reduction ratio

References 1 Milk and milk products in the E.U.

2

3

4

5

6

European Commission. Directorate for Agriculture (2006) http://ec.europe.eu/ agriculture/ Walstra, P., Geurts, T.J., Noomen, A., Jellema, A. et al. (1999) Dairy Technology: Principles of Milk Properties and Processes, Marcel Dekker, New York. Brans, G., Schroen, K., Van der Sman, R.G.M. et al. (2004) Membrane fractionation of milk: state of art and challenges. J. Membrane Sci., 243, 263–272. Cayot, P. and Lorient, D. (1997) Structure-function relationships of whey proteins, Food Proteins and Their Applications (eds. S. Damodaran and A. Paraf), Marcel Dekker, New York, pp. 225–256. Smit, G. (2003) Dairy Processing: Improving Quality, Woodhead Publishing Limited, Cambridge, England, ISBN: 1 85573 7007 8. Huffman, L. and Harper, J. (1999) Maximizing the value of milk through

7

8

9

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11

12

separation technologies. J. Dairy Sci., 82, 2238–2244. Smithers, G.F., Ballard, J., Copeland, A.D., de Silva, K.J., Dionysius, D.A. et al. (1996) New opportunities from de isolation and utilization of whey proteins. J. Dairy Sci., 79, 1454–1459. Perez-Gago, M. and Krochta, J. (2002) Formation and Properties of whey proteins ﬁlms and coatings, Protein-based Films and Coatings (ed. A. Gennadios), CRC Press, USA, ISBN: 1-587-16107-99. Merin, U. and Dauﬁn, G. (1990) Crossﬂow ﬁltration in the dairy industry: state of the art. Lait, 70, 281–291. Saboya, J.L. and Maubois, J.L. (2000) Current development of microﬁltration technology in the dairy industry. Lait, 80, 541–553. Dauﬁn, G., Radenac, J.F., Gesan, G. et al. (1993) A novel rig design for ultra- and microﬁltration experiments. Separ. Sci. Technol., 28, 2635–2642. Gesan-Guiziou, G., Boyaval, E., and Dauﬁn, G. (1999) Critical stability

References

13

14

15

16

17

18

19

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21

conditions in crossﬂow microﬁltration of skimmed milk: transition to irreversible deposition. J. Membrane Sci., 158, 211–232. Pouliot, M., Pouliot, Y., and Britten, M. (1996) On the conventional cross-ﬂow ﬁltration of milk for the production of native caseinate. Int. Dairy J., 6, 105–111. Vadi, P.K. and Rizvi, S.S.H. (2001) Experimental evaluation of a uniform transmembrane pressure crossﬂow microﬁltration unit for the concentration of micellar casein from skim milk. J. Membrane Sci., 189, 69–82. Le Berre, O. and Dauﬁn, G. (1996) Skimmilk crossﬂow ﬁltration performance versus permeation ﬂux to wall shear rate. J. Membrane Sci., 117, 261–270. Samuelsson, G., Dejlmek, P., Tragardh, G., and Paulsson, M. (1997) Minimizing whey protein retention in crossﬂow microﬁltration of skim milk. Int. Dairy J., 7, 237–242. Lawrence, N.D., Kentish, S.E., Connor, A.J., Barber, A.R., and Stevens, G.W. (2008) Microﬁltration of skim milk using polymeric membranes for casein concentrate manufacture. Sep. Purif. Technol., 60, 237–244. Nelson, B.K. and Barbano, D.M. (2005) A microﬁltration process to maximize removal of serum proteins from skim milk before cheese making. J. Dairy Sci., 88, 1891–1990. Zulewska, J., Newbold, M., and Barbano, D.M. (2009) Efﬁciency of serum protein removal from skim milk with ceramic and polymeric membranes at 50 C. J. Dairy Sci., 92, 1361–1377. Dauﬁn, G., Pierre, A., Garem, A., Goudedranche, H., and Gesan-Guiziou, G. (1998) Industrie laitiere, Chapter 7, Les Separations Par Membrane Dans Les Procedes De Lindustrie Alimentaire (eds. G. Dauﬁn, F. Rene and P. Aimar), Tec Doc, Lavoisier, Paris, pp. 282–371. Ernstrom, C.A., Sutherland, B.J., and Jameson, G.W. (1980) Cheese base for processing a high yield product from whole milk by ultraﬁltration. J. Dairy Sci., 63, 228–234.

22 Labbe, J.P., Quemerais, A., Michel, F.,

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and Dauﬁn, G. (1990) Fouling of inorganic membranes during whey ultraﬁltration: analytical methodology. J. Membrane Sci., 51, 299–307. Zydney, A.L. (1998) Protein separation using membrane ﬁltration: new opportunities for whey fractionation. Int. Dairy J., 8, 243–250. Clarke, T.E. and Heath, C. (1997) Ultraﬁltration of skim milk in ﬂat-plate and spiral-wound modules. J. Food Eng., 33, 373–383. Yan, S.H., Hill, C.G., and Amundson, C.H. (1979) Ultraﬁltration of whole milk. J. Dairy Sci., 62, 23–40. Cheang, B. and Zydney, A.L. (2003) Separation of a-Lactalbumin and b-Lactoglobulin using membrane ultraﬁltration. Biotechnol. Bioeng., 83, 201–209. Cheang, B. and Zydney, A.L. (2004) A two-stage ultraﬁltration process for fractionation of whey protein isolate. J. Membrane Sci., 231, 159–167. Muller, A., Chaufer, B., Merin, U., and Dauﬁn, G. (2003) Prepuriﬁcation of a-lactalbumin with ultraﬁltration ceramic membranes from casein whey: study of operating conditions. Lait, 83, 111–129. Almécija, M.C., Ibanez, R., Guadix, A., and Guadix, E.M. (2007) Effect of pH on the fractionation of whey proteins with a ceramic ultraﬁltration membrane. J. Membrane Sci., 288, 28–35. Bramaud, C., Aimar, P., and Dauﬁn, G. (1997) Optimization of a whey protein fractionation process based on the selective precipitation of a-lactalbumin. Lait, 77, 411–423. Lucas, D., Rabiller-Baudry, M., Millisime, L., Chaufer, B., and Dauﬁn, G. (1998) Extraction of a-lactalbumin from wheyprotein concentrate with modiﬁed inorganic membranes. J. Membrane Sci., 148, 1–12. Kelly, J. and Kelly, P.M. (1995) Nanoﬁltration of whey: quality, environmental and economic aspects. J. Soc. Dairy Technol., 48, 20–25. Dauﬁn, G., Escudier, J.-P., Carrere, H., Berot, S., Lillaudeau, L., and Decloux, M. (2001) Recent and emerging applications

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of membrane process in the food and dairy Industry. Trans. Inst. Chem. Eng., 79, 89–102. Balannec, B., Gésan-Guiziou, G., Chaufer, B., Rabiller-Baudry, M., and Dauﬁn, G. (2002) Treatment of dairy process waters by membrane operations for water reuse and milk constituents concentration. Desalination, 147, 89–94. Balannec, B., Vourch, M., Rabiller-Baudry, M., and Chaufer, B. (2005) Comparative study of different nanoﬁltration and reverse osmosis membranes for dairy efﬂuent treatment by dead end ﬁltration. Sep. Purif. Technol., 42, 195–200. Vourch, M., Balannec, B., Chaufer, B., and Dorange, G. (2008) Treatment of dairy industry wastewater by reverse osmosis for water reuse. Desalination, 219, 190–202. Jaffrin, M.Y. (2008) Dynamic shearenhanced membrane ﬁltration: a review of rotating disks, rotating membranes and vibrating systems. J. Membrane Sci., 324, 7–25. Kroner, K.H. and Nissingen, V. (1988) Dynamic ﬁltration of microbial suspensions using an axially rotating ﬁlter. J. Membrane Sci., 36, 85–100. M€antt€ari, M., Vitikko, K., and Nystrom, M. (2006) Nanoﬁltration of biologically treated efﬂuents from the pulp and paper industry. J. Membrane Sci., 272, 152–160. Kaiser, B. (2004) Dynamic crossﬂow ﬁltration with rotating membranes. Proceed. Euromembrane Congress, Sept 28-Oct 1, Hamburg, All, p. 620. Al-Akoum, O., Jaffrin, M.Y., Ding, L.H., Paullier, P., and Vanhoutte, C. (2002) An hydrodynamic investigation of microﬁltration and ultraﬁltration in a vibrating membrane module. J. Membrane Sci., 197, 37–52. Zouboulis, A.I. and Petala, M.D. (2008) Performance of VSEP vibratory membrane ﬁltration system during the treatment of landﬁll leachates. Desalination, 222, 165–175. Postlethwaite, J., Lamping, S.R., Leach, G.C., Hurwitz, M.F., and Lye, G.J. (2004) Flux and transmission characteristics of a vibrating microﬁltration system operated

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at high biomass loading. J. Membrane Sci., 228, 89–101. Al-Akoum, O., Ding, L.H., and Jaffrin, M.Y. (2002) Microﬁltration and ultraﬁltration of UHT skim milk with a vibrating membrane module. Sep. Purif. Technol., 28, 219–234. Al-Akoum, O., Ding, L.H., ChotardGhodsnia, R., Jaffrin, M.Y., and GesanGuiziou, G. (2002) Casein micelles separation from skimmed milk using a VSEP dynamic ﬁltration module. Desalination, 144, 325–330. Espina, V.S., Jaffrin, M.Y., Frappart, M., and Ding, L.H. (2008) Separation of casein micelles from whey proteins by high shear microﬁltration of skim milk using rotating ceramic membranes and organic membranes in a rotating disk module. J. Membrane Sci., 325, 872–879. Blatt, W.F., Dravid, A., Michaels, A.S., and Nelson, L. (1970) Solute polarization and cake formation in membrane ultraﬁltration: causes, consequences and control techniques, Membrane Science and Technology (ed. J.E. Flinn), Plenum Press, New York, N.Y., pp. 47–97. Jaffrin, M.Y., Ding, L.H., Akoum, O., and Brou, A. (2004) A hydrodynamic comparison between rotating disk and vibratory dynamic ﬁltration systems. J. Membrane Sci., 242, 155–167. Ding, L.H., Akoum, O., Abraham, A., and Jaffrin, M.Y. (2003) High shear skim milk ultraﬁltration using rotating disk ﬁltration systems. AIChE, 49, 2433–2441. Akoum, O., Jaffrin, M.Y., and Ding, L.H. (2005) Concentration of total milk proteins by high shear ultraﬁltration in a vibrating membrane module. J. Membrane Sci., 247, 211–220. Espina, V., Jaffrin, M.Y., and Ding, L.H. (2009) Extraction and separation of a-Lactalbumin and b-Lactoglobulin from skim milk by microﬁltration and ultraﬁltration at high shear rates: a feasibility study. Sep. Science Tech., 44, 3832–3853. Gesan-Guiziou, G., Dauﬁn, G., Timmer, M., Allersma, D., and van der Horst, C. (1999) Process steps for the preparation of puriﬁed fractions of a-lactalbumin from

References whey protein concentrates. J. Dairy Res., 66, 225–236. 53 Bhattacharjee, S., Bhattacharjee, C., and Datta, S. (2006) Studies on the fractionation of b-lactoglobulin from casein whey using ultraﬁltration and ionexchange membrane chromatography. J. Membrane Sci., 275, 151–150. 54 Akoum, O., Jaffrin, M.Y., Ding, L.H., and Frappart, M. (2004) Treatment of

dairy process waters using a vibrating ﬁltration system and NF and RO membranes. J. Membrane Sci., 235, 111–122. 55 Frappart, M., Akoum, O., Ding, L.H., and Jaffrin, M.Y. (2006) Treatment of dairy process waters modelled by diluted milk using dynamic nanoﬁltration with a rotating disk module. J. Membrane Sci., 282, 465–472.

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4 Electrodialysis in the Food Industry Jamie Hestekin, Thang Ho, and Thomas Potts 4.1 Introduction

During the last century, electrodialysis developed from a laboratory curiosity to a powerful tool that is applicable in a wide variety of industrial applications. Of special interest is the application of the variations of electrodialysis to difﬁcult separations found in certain food industries. Electrodialysis is often the separation tool of choice in the dairy, wine, and juice and sugar industries. The ﬁrst experiments with ion-exchange membranes were performed in the early 1890s by Ostwald, and opened many opportunities for membrane-separation technology [1]. The concepts of membrane potential and the Donnan exclusion phenomenon were developed a few years later [2]. The concept of electrodialysis was introduced by Manegold and Kalauch [3] in 1940. They arranged cationic and anionic ion-exchange membranes to separate ions from water. In that same year, Meyer and Strauss expanded this to assemble many such membrane pairs into a multicell arrangement between a single pair of electrodes [4]. Using this arrangement with the newly developed polymer membranes, electrodialysis quickly became the technology of choice for commercial desalination plants. A variation of this technology, electrodialysis reversal (EDR), was developed in the 1970s to address certain problems characterizing traditional electrodialysis [5, 6]. EDR exhibited lower operating costs, especially in membrane-system maintenance, and replaced traditional electrodialysis in desalination applications. The lower operating costs associated with EDR also facilitated extension of electrodialysis to other commercial separation problems. The development of bipolar membranes in 1977 and perﬂouro-based membranes in 1979 further expanded the applicability of electrodialysis technology to a wide variety of industrial separations [7]. Electrodialysis in its many variations has become not only become a technology for desalination, but also a viable and cost-effective solution for separation problems throughout the dairy, juice, and wine industries. In this chapter, electrodialysis theory and applications as applied to the food industry will be discussed.

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4.2 Technology Overview 4.2.1 Principle of the Electrodialysis Process

Electrodialysis is the separation of ionic materials under the inﬂuence of an electric ﬁeld in a system comprised of ion-exchange membranes arranged to make ﬂow compartments called cells. The three types of ion-exchange membranes commonly employed in electrodialysis systems are anion-exchange membranes, cation-exchange membranes, and bipolar membranes. Anion-exchange membranes are membranes that allow anions to pass through (permeate) but do not allow cations to permeate. Cation-exchange membranes are membranes that permeate cations but not anions. An example of the action of a cation-exchange membrane is shown in Figure 4.1. Bipolar membranes (Figure 4.2) are the lamination of a cation-exchange membrane and anion-exchange membrane. They do not allow ions of either charge to permeate all the way through the membrane. Bipolar membranes are primarily used to produce acids and bases by electrolysis of salt solutions [8, 9]. Electrodialysis utilizes the chemistry of the membranes and an electrical potential to remove ions

Figure 4.1 Cation-exchange membrane action.

4.2 Technology Overview

Figure 4.2 Bipolar membrane.

from solutions. A typical electrodialysis setup consists of cation- and anion-exchange membranes arranged alternately and separated slightly with some form of spacer. Each pair of cation- and anion-exchange membranes is called a cell pair. A series of cell pairs put together between two electrodes is called a stack (Figure 4.3). The number of cell pairs used in a stack depends on the requirements of the speciﬁc application and may reach as many as 500 cell pairs [10, 11]. Electrodialysis systems transport ions through the sum of two different driving forces: ion concentration gradients and electric potential gradients. The forces generated by the electrical potential gradients in electrodialysis systems are usually much larger than the forces generated by ion concentration gradients. In electrodialysis, cations and anions are transported in opposite directions; however, the fraction of electrical current carried by cations and anions is not necessarily equal. The fraction of total electric current carried by either cations or anions is called the transport number of that type of ion. The sum of all transport numbers is one in electrodialysis systems. The cation transport number is a function of the velocity of the cations (u) in the externally applied electric ﬁeld and is shown in Equation 4.1. The anion transport number is a function of the velocity of the anions measured in the same direction (v) and is shown in Equation 4.2 [8–10] tþ ¼

u uþv

ð4:1Þ

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Figure 4.3 Electrodialysis stack.

t ¼

v uþv

ð4:2Þ

where t þ is the cation transport number and t is the anion transport number. The transport numbers of ions are different for different ionic species and reﬂect the different sizes and charges of the ions. In electrodialysis or other membraneseparation systems, the transport number of ions having the same charge as the charge of the ion-exchange membranes approaches zero. The transport number approaches one for ions having a charge opposite of the charged group of the membrane. A transport number close to zero means the membrane does not allow ions to permeate, while a transport number close to one means the ions can pass easily through the membrane. A difference in transport number allows separations to be achieved with ion-exchange membranes. Manipulation of concentrations of salts and ions in electrodialysis systems can help achieve the separation of interest. For instance, consider the separation of cations and anions by a cation-exchange membrane. The concentration of cation in þ the membrane is deﬁned as cðmÞ , the concentration of anion in the membrane is deﬁned as cðmÞ , the concentration of ﬁxed-charged group inside the membrane is cR (m), and the concentration of salt in the solution is c(s). The mathematical expression of cation transport inside the membrane is shown in Equation 4.3 [12–14]: þ cðmÞ

cðmÞ

¼

1 cR ðmÞ 2 k cðsÞ

ð4:3Þ

where k is an equilibrium constant. Equation 4.3 shows that as the concentration of the salt in the feed solution increases, the ratio of cation concentration to anion

4.2 Technology Overview

concentration in the membrane decreases. This means the cation-exchange membrane becomes partially permeable to anions and the overall membrane separation becomes poorer. More details about ion transport in membrane separation and electrodialysis theory can be found from a variety of sources [6, 10, 12–24]. 4.2.2 System Design 4.2.2.1 Concentration Polarization, Limiting Current Density, Current Utilization, and Power Consumption The performance of electrodialysis systems is determined by several factors. In a typical electrodialysis stack, ions migrate through a membrane but do not permeate the next membrane in the stack. For instance, cations migrate through a cationexchange membrane but do not pass through the next anion-exchange membrane. As a result, the salt concentration increases in those compartments where the ions cannot exit, while the other compartments are continuously depleted of salts. In the typical stack arrangement, every other compartment becomes more concentrated, while the alternating compartments become less concentrated. The chambers where the salt concentration increases are called the concentrate compartments, and the chambers where the salt concentration decreases are called the diluate compartments. In a typical electrodialysis system, the efﬁciency of the system is usually dictated by the electrical resistance of the diluate compartments. This resistance is high because the low ion concentration in the compartment does not support electrical current conduction. The formation of a low ion concentration boundary layer at the membrane surface also lowers the separation efﬁciency. The difference between the bulk solution ion concentration and the low ion concentration layer at the surface of the membrane is called concentration polarization. Performance of an electrodialysis system is limited by concentration polarization. As the ion concentration adjacent to the membrane decreases, the electrical potential must be increased to maintain the same ion ﬂux across the membrane. As ions transport through the membrane, the concentration of ions next to the membrane surface becomes smaller. Ion transport is thus limited by the depleted layer at the membrane surface. The energy consumption per ion transported increases signiﬁcantly when concentration polarization occurs. When the ion concentration at the membrane surface approaches zero, the transport rate of ions through the membrane becomes the transport rate through the boundary layer. The electric current per membrane area through the electrodialysis system at this point is called the limiting current density. Increasing the current above the limiting current density will not increase the ion transport across the membrane and this is thus wasted power. The excess power most commonly goes to the dissociation of water into hydrogen and oxygen [25, 26]. The energy used in water dissociation is wasted power, and commercial systems are operated below the limiting current density. Concentration polarization can be reduced to some extent by thoroughly mixing the salt solution in each compartment but will never be eliminated in electrodialysis systems. The adverse effects of concentration polarization signiﬁcantly reduce the

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performance of electrodialysis systems. If the concentration of ions in the compartment becomes very low, the internal resistance of the compartment becomes very high, and limiting current density can occur even without concentration polarization. The limiting current density is highly speciﬁc for each electrodialysis system and can best be determined through experimentation. The limiting current density can be determined by measuring the total resistance of the stack and the pH of the diluate chamber as a function of current density. When the pH is plotted versus the reciprocal of current, a sharp decrease in pH is noted when the limiting current is reached. Similarly, when the total resistance of a stack is plotted versus the reciprocal of the current, a minimum is obtained, indicating the limiting current density. Determination of the limiting current density is rather difﬁcult in industrial-scale electrodialysis systems and it is usually approximated in practice [9, 13, 18, 27]. The performance of an electrodialysis system is usually evaluated by the energy consumption required to perform a separation. The energy consumption (E ) is a function of the voltage (V ) applied across the system and the current (I ) through the stack as shown in Equation 4.4 E ¼ IV

ð4:4Þ

Another way to evaluate system performance is to calculate the current utilization. Current utilization is the ratio of theoretical current required to transport charges across the membrane to the actual operating current. The theoretical current is a function of the valence (z) of the ion, the change in concentration of ions (DC), Faradays constant (F), and the solution ﬂow rate (Q) as shown in Equation 4.5 Itheor ¼ zDCFQ

ð4:5Þ

Current utilization is always less than 100% (usually greater than 90%), this is due to many factors and the discussion of these factors is beyond the scope of this discussion. However, this is a simple calculation that allows one to determine how well an ED system is operating. Details of concentration polarization, limiting current density, power consumption, and current utilization can be found in many papers [6, 12, 13, 20, 21, 23, 28, 30–32]. 4.2.2.2 System Design and Cost Analysis Many factors must be incorporated into the design and cost analysis of an electrodialysis system. Some general comments relating to electrodialysis design and cost analysis are made below. Electrodialysis can be a single-stage or multistage process, depending on the application. For either arrangement, a typical electrodialysis system consists of ﬁve components: a feed pretreatment system, a membrane stack, a power supply, a control system and a pumping system. The feed pretreatment is necessary to prevent membrane fouling by particle deposition on the membrane surface. The pretreatment process needed depends on the feed quality and is usually a combination of microﬁltration or ultraﬁltration and pH adjustment or addition of antiscaling chemicals. For instance, in the wine industry, the pretreatment of wine after

4.2 Technology Overview

fermentation but before electrodialysis often includes centrifugation and reverse osmosis to remove solid particles [33, 34]. In the production of lactic acid from whey, pretreatment often consists of pH adjustment and microﬁltration [35–37]. The membrane stacks in electrodialysis systems consist of up to 500 membrane cell pairs with an active membrane area of 1–2 m2 per cell pair [13, 33, 38–41]. Between the membranes of each compartment, there is a spacer to evenly distribute the process ﬂow. The two most common spacers used are tortuous path (Figure 4.4a) and sheet ﬂow (Figure 4.4b). Tortuous path membranes must be thicker and sturdier than sheet-ﬂow membranes since there is no additional spacer between the membranes in the tortuous path system. The choice of spacer is often dictated by

Figure 4.4

(a) Tortuous path stack spacer and (b) Sheet flow path stack spacer.

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preference of the membrane vendor. The length of the ﬂow path between membranes is designed as short as practically possible to decrease ﬂuid ﬂow resistance and pressure drop. Cleaning of an ED stack in food applications is similar to other membrane process with various acid, bases, and sanitization rinsing steps. However, since chlorine is seldom used because of membrane degradation, membrane cleaning and replacement often needs to take place more often in food applications. In theory, the electrodialysis stack can be disassembled and the membranes cleaned and replaced on site when they become heavily soiled. However, in situ cleaning is performed infrequently because it is difﬁcult to reassemble an electrodialysis stack without introducing leakage. Membrane cleaning is usually done by the vendor and is performed once or twice a year, depending on the type of application [42–45]. The power supply and process control units of an electrodialysis system comprise a large portion of the capital cost of an electrodialysis system. The electric current used in electrodialysis systems is usually direct current (DC) rather than alternating current (AC). Electrodialysis systems operate at high voltage and high current, which requires stringent precautions to ensure safe operation. Such precautions include, but are not limited to, good electrical insulation around the system and periodic checks for corroded parts [10, 11, 17, 28, 46]. The last component comprising electrodialysis systems is the pumping system. Typical pressure drops in stacks vary from 15 to 30 psi for a sheet ﬂow path cell and from 70–90 psi for a tortuous path cell [13, 18, 47]. Depending on the application, interstage pumps might be necessary for the stack. In a multistage electrodialysis system, power consumption by the pumping system is a large fraction of the plant operating cost. This power consumption fraction increases as the concentration of feed or diluate decreases, because less power is required for separations, and more power is required for mixing. The total cost of ownership for electrodialysis systems consists of many capital and operating costs. The depreciable capital cost items in electrodialysis systems are membrane stacks, pumps, electrical equipment, and control units. The capital investment required for electrodialysis plants is dictated by the total number of ions that must be removed from the feed solution. The lifetime of membranes is usually assumed to be 5 years and the lifetime of other equipment is usually assumed to be 10 years. With these membrane lifetimes, the operating cost of electrodialysis plants is dominated by energy consumption (>90% of total operating cost) [10, 11, 20, 21, 23, 24, 48, 49]. The energy cost can be calculated from the energy required for the separation process and the energy for the pumping systems. Details of the economic analysis of electrodialysis systems can be found in the following references [6, 10–12, 20, 23, 24, 29, 32, 50].

4.3 Electrodialysis Applications in the Food Industry

The use of membranes in the food industry has increased steadily for the past 25 years. In 1988, the total annual sale of membranes and membrane modules for

4.3 Electrodialysis Applications in the Food Industry

food applications was estimated at about 160 million USD, or about 15% of the total annual sales of this industry. By 2001, the total annual sales increased to 400 million USD, or by 7.5% per year [51]. Membrane technologies used in the food-processing industry include reverse osmosis, ultraﬁltration, microﬁltration, and electrodialysis. Electrodialysis systems annual sales account for about 10% of the total membranes systems sold [9, 41, 43]. The main applications of electrodialysis are in dairy (40%) and beverages (wine, beer, fruit juices, etc.) [9, 52]. Additionally, there are emerging electrodialysis processes for the treatment and transformation of raw agricultural products into safe and well-accepted food products. Pertinent characteristics of electrodialysis systems adopted by the food industry are [13, 53–57] . . .

improvement of process performance and food quality in preparation of traditional food products; innovation of processes and products aimed at satisfying evolving food requirements related to nutrition and health; meeting the demands of changing regulations related to waste and waste treatment in food processes.

Electrodialysis gives the food industry three advantages as compared to competing technologies: increased food safety, economic competitiveness, and environmental friendliness. Current applications of electrodialysis in the juice, wine, and dairy industries highlight the innovation and diversity of electrodialysis in food processing. 4.3.1 Wine Industry

Electrodialysis is commonly used by the wine industry to remove tartrate salts from wine before bottling. Tartrate salts have a tendency to precipitate during storage and the precipitates decrease the quality of wine. A block diagram of a process for making wine from grape must with integrated membrane technology is shown in Figure 4.5

Figure 4.5 Process flow diagram for making wine from grape must.

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[22, 28, 34, 54, 58, 59]. Grape must is ﬁrst centrifuged to remove solid particles and then passed through either an ultraﬁltration or microﬁltration unit to remove microorganisms. A portion of the sterile must is then passed through a reverse osmosis unit to increase its sugar concentration. The concentrated must is then blended with the remaining must to achieve a desired sugar level before sending the must concentrate to the fermentation step. Yeast starter is added to the fermenter to convert the concentrated must to wine. The product from the fermenter is either centrifuged or ﬁltered to remove the lees. In the last step, the wine product is either treated with electrodialysis or chilled and ﬁltered by microﬁltration to reduce the levels of tartrate salts. In Figure 4.5, electrodialysis is used to prevent the precipitation of tartrate salts in the wine product. Electrodialysis is also sometimes used before fermentation to stabilize the ﬁnal product. Other salts can also be problematic in wine production. These salts are naturally present in the grape must and can be precipitated as potassium bitartrate, potassium bimalate, potassium tartrate, calcium tartrate, calcium malate, calcium succinate, and calcium oxalate [22, 33, 38, 50, 56, 58, 60, 61]. Electrodialysis removes excess salts from wines or grape juices. The amount of ions needed to be removed from the solution is dependent upon the type of wine, grapes, and type of vineyards. It is difﬁcult to generalize the optimal amount of these ions at the various stages of the wine production. Some studies of red wine suggest that the amount of potassium should be reduced to a level of 100 to 450 mg/l [28, 34, 45, 50, 62–64], depending on the type of wine, while other studies suggest a 10% decrease in the concentration of potassium ion is enough to stabilize white wine [28, 44, 50, 56, 61]. The removal of cations increases the acidity of the wine or grapes and decreases the alcohol content of the wine [42, 47, 50, 56, 65]. Moreover, a 10% sugar loss has been reported during demineralization of must using electrodialysis [28, 42, 43, 56, 65]. The presence of sulfur dioxide helps to stabilize the wine products from spoiling due to microorganisms. However, electrodialysis systems extract HSO3 at a very high rate. Approximately 50–80% of the total SO2 is eliminated from musts containing up to 850 mg l1 of SO2. For wines with a low SO2 concentration (100 ppm), only 20% of the SO2 is extracted [44, 61, 65]. Sulfonic components of the must are not affected by electrodialysis and their concentration remains constant through electrodialysis [42, 44, 56, 61]. Other organic acids such as malic acid are removed at the same rate as tartaric acid [33, 34, 54]. In the presence of high tannin and anthocyanin concentrations, typical in red wines, potassium and tartaric ion removal are decreased. Electrodialysis is also used in deacidiﬁcation and acidiﬁcation of grape musts and wine to harmonize the wine by adjusting the sugar content, acid content, and pH. A special conﬁguration of electrodialysis ion substitution is used for this purpose. In ion substitution, the electrodialysis system works as an ion exchanger. Two anionic membranes are put together to create a cell pair instead of the more usual cell cation and anion membranes. Three different compartments are formed with three ﬂow streams: acceptor, donor, and product [38]. The donor solution donates ions into the product stream while the acceptor stream receives ions from the product stream. The product stream receives ions from the donor stream and delivers different ions to

4.3 Electrodialysis Applications in the Food Industry

the acceptor stream. For instance, when NaOH or KOH is used as the donor stream, OH will replace the anion group in the feed or product stream and make the stream less acidic. It has been reported that the acid concentration in wine can be reduced from 7.0 to 3.7 g/l using deacidiﬁcation electrodialysis [22, 28, 33, 42, 45, 62, 63, 66, 67]. For acidiﬁcation of grape must and wines, two cation membranes are used to form the three compartments and streams. The donor stream in this case is an acidic solution. The pH of wine can be adjusted from 4.5 to 3 using an acidiﬁcation electrodialysis process [28, 48, 53, 66–70]. It has been reported that the current efﬁciency for electrodialysis systems in grape juice and wine stabilization is between 65–75% depending on the quality of the feed [45, 48, 53, 67–69]. As addressed in the previous section, power consumption is directly proportional to the current density. For grape must and wine treatments, a current density at 100 A m2 leads to an energy consumption of about 5.0 kWh kg1 of K þ removed [45, 68, 71]. The energy consumption during the electrodialysis process for concentrating must is between 3 and 4.4 kWh m3 of treated must at the beginning of the concentration process. The energy consumption increases as the feed becomes more dilute and becomes 17 kWh m3 during the last stages of the process [67]. It has been reported that the typical cost of electrodialysis systems sized for a production rate of 10 million gallons per year is about $400 000 with an operating cost of $0.01/l of wine. For vineyards with low capacity (less than 4500 m3 year1), an electrodialysis system can be rented for less than $0.10 per bottle of wine [28, 34, 38, 61]. 4.3.2 Juice and Sugar Industry

The two primary applications of electrodialysis in the juice and sugar industry are deacidiﬁcation and demineralization. The juice extracts from orange, grape, pineapple, and lemon are highly acidic. Acid concentrations of 1.0–1.2% in orange, grape, and pineapple juices interfere with utilization of these juices in single-strength or concentrated forms [8, 73]. About 15–25% of the pineapple juice obtained as byproduct in the pineapple canning industry is not suitable for production of singlestrength or concentrated juice due to high acidity [48, 67, 69, 70, 74–76]. The sourness or sweetness in the juices is related to the ratio of soluble solids (sugars) to acids in the juice. The concentration of sugars in the fruits remains constant during the growing season but the concentration of citric acid increases during the fall and winter months. In the juice industry, the ratio of soluble solids to acid in the juice is called the Brix/acid ratio. A Brix/acid ratio of less than about 12 is undesirably sour for orange juice; a sweet orange juice has a Brix/acid ratio of 13.5–14.5 [42, 45, 67, 72, 77, 78]. A Brix/acid ratio in grapefruit juice of less than 9 is too sour for consumption; an acceptable Brix/acid ratio for grapefruit juice is about 10 to 11 [8, 42, 45, 52]. The Brix/acid ratio of sour juice can be increased to a desirable range by blending the sour juice with high naturally sweet juices that have been saved from the previous harvest season. However, there is usually not enough natural sweet juice available for blending with sour juices. There is also signiﬁcant cost involved with storage of

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naturally sweet juices when using the blending option. Another method to increase the Brix/acid ratio is through sugar addition. This procedure suffers from the high cost of sugar and blending equipment. Further, legal requirements mandate that canned or frozen juices to which sugar is added must be labeled Sugar Added. The juice industry has found that sugar-added products command a lower price than products for which the label is not needed. Moreover, if the juice is exceptionally sour, the quantity of sugar needed to raise the Brix/acid ratio to acceptable levels may cause a syrupy consistency of the juice. Using alkaline materials to neutralize the acid in the juice can also increase the Brix/acid ratio. However, this method causes unacceptable changes in ﬂavor and/or formation of undesirable precipitates. Deacidiﬁcation using electrodialysis eliminates all of these storage and legal problems. The electrodialysis stack used for deacidiﬁcation of juices consists of two anion-exchange membranes. The stack formed from these cells consists of alternating diluate compartments (juice compartments) and concentrate compartments (alkali compartments). In this conﬁguration, only anions pass through membranes and the net effect is the extraction of anions from the juice and their replacement by OH ions from the alkali compartment. The voltage potential is periodically reversed without interchanging the two streams (this technique is referred to as electrodialysis reversal) to prevent colloids and solids from depositing on the surface of the membrane. The energy requirement for juice deacidiﬁcation varies between 0.02 and 0.1 kWh/equiv., which is between 6 and 10 kWh m1. The current efﬁciency for an electrodialysis system in the deacidiﬁcation of fruit juice is from 52 to 90% depending on the quality of the juice [8, 47, 52, 73, 79–81]. Cloudy or unclariﬁed apple juice is in high demand because of its high content of dietary ﬁber and important nutrients. However, it is difﬁcult to produce superiorquality cloudy juice. Cloudy apple juice is very sensitive to enzymatic browning because it contains high quantities of polyphenols and polyphenol oxidase (PPO). The enzymatic browning reaction is catalysed by PPO and coverts polyphenol to o-quinones, which then polymerize to form complex dark pigments. Therefore, the composition of the apple juice changes as the reaction occurs. Temporarily lowering the pH of apple juice to 2.0 and then readjusting the pH to normal values will irreversibly inhibit PPO activity and stabilize the juice color [82–84]. The previous approach to this process was to use hydrochloric acid and caustic soda to adjust the pH of the juice; however, this treatment results in the formation of salts that adversely affect the ﬂavor of the apple juice. Acidiﬁcation of apple juices by bipolar membrane electrodialysis avoids the formation of ﬂavor-degrading salts. As discussed previously, bipolar membranes are membranes having the characteristics of both cationic and anionic membranes. Bipolar membranes are used to produce acids and bases by electrodialysis. With this unique characteristic, bipolar membrane electrodialysis is a perfect tool to adjust the pH in cloudy apple juices for enzyme inhibition, as shown in Figure 4.6. In this electrodialysis system, potassium chloride solution is used as the concentrate solution. Potassium ions in the juice migrate across the cationic membrane into the concentrate compartment and are replaced by hydrogen ions formed at the bipolar membrane. Using this conﬁguration, the pH of apple juice can be lowered from 3.5 to 2, and the enzymatic activity decreases signiﬁcantly.

4.3 Electrodialysis Applications in the Food Industry

Figure 4.6 pH Adjustment of apple juice.

The energy consumption for this process ranges from 20–97 kWh/m3 of juice. The current efﬁciency based on the amount of potassium removed from apple juice solution ranges from 60–90%, depending on the quality of the apple juice [53, 66, 69, 70, 82–84]. Another important application of electrodialysis in juice and sugar industries is to reduce the mineral content (demineralize) of sugar sirup. It has been known for more than 40 years that alkali metal cations are highly melassigenic; they hold sugar in the molasses and prevent it from being recovered as crystalline white sugar. Many authors quantify the melassigenic effect of the alkali and alkaline-earth ions. The affects of these ions decreases in the order K > Na > Ca > Mg, with the potassium and sodium ions much more melassigenic than magnesium ions. The raw juice of sugar beet or sugar cane contains up to 3.5% ionized materials [57, 77, 85–91]. These ionized impurities inhibit the crystallization of sugar and cause scaling of the tubes in the evaporators. It has been shown that if the ions are removed from the juice, about 5% more sugar is recovered from each ton of cane or beets, and scaling in the evaporator tubes is reduced [11, 42, 89, 92]. Several technologies have been employed in the sugar industry to remove melassigenic ions: ion-exchange resins, synthetic adsorbents, coagulants and membranes. However, these technologies are costly and have a short lifetime in high sugar content solutions. Electrodialysis systems containing cation and anion membranes in alternating order (the usual conﬁguration) are used in the demineralization of sugar sirup. A problem encountered with the use of electrodialysis in ion removal from sugar juices is a high fouling potential in systems that are not properly cleaned. Fouling in these systems is caused by negatively charged organic materials in the sugar juice. These materials deposit on the surface of the membranes and increase the resistance of the membranes, which in turn decreases the current efﬁciency. Membrane fouling reduces the production rate and increases the energy consumption costs. A properly working and foulingfree electrodialysis system requires about 1.10 kWh kg1 of salt removed for juice applications [80, 85, 93]. The energy savings by using electrodialysis is 440 kWh ton1

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of sugar produced [11, 57, 85]. The typical current efﬁciency of electrodialysis systems in this application is between 40–45%, depending on the sugar content of the sirup [17, 42, 44, 89, 92]. Recent studies indicate that electrodialysis use will recover sugar and potassium from blackstrap molasses. Blackstrap molasses is the liquid left from the crystallization of sugar. Blackstrap molasses contains about 55% sucrose as invert sugar, 10% ash, 5–10% nonsugar organic materials, and 18–25% water [85, 86, 89, 92]. The inorganic and nonsugar organic materials inhibit the crystallization of the sucrose. Blackstrap molasses is usually sold at low price for cattle feed or for alcohol production. The ash in blackstrap is mostly potassium compounds and is valuable for fertilizer production. At current prices, blackstrap molasses is much lower in value than the sucrose and potassium contained in the solution. Electrodialysis for the recovery of potassium and sugar from blackstrap molasses has the potential to become a high value added process for the sugar industry. Research and economic investments in this emerging technology continues. 4.3.3 Dairy Industry

Electrodialysis is used to demineralize and acidify whey in the dairy industry. Whey is the ﬂuid by-product in cheese manufacture. In the United States, cheese manufacturers produce about 25 billion pounds of whey yearly. The whey contains highly nutritious materials: 12% protein, 1% fat, 70–75% lactose, 8–10% ash, and 0.1–1% lactic acid (based on dry weight). Whey is a good source of protein, milk, sugar, and vitamins; however, its high ash content makes its unsuitable as human food [35, 36, 94, 95]. There are two different types of whey: one from the curd in cheese making and one from casein production [43–45, 47, 72, 96, 97]. The compositions of the two types are similar. In spite of its high ash content, a portion of whey is dried and sold at low prices as an additive in animal feed. Ultraﬁltration is used to recover protein from whey. The product of using this technology is high-grade protein suitable as human food. However, the large amount of lactose in the ultraﬁltration permeate still results in a serious waste-disposal problem. The worldwide capacity for whey desalting by electrodialysis is about 100 000 tons per year of 90% demineralized whey powder from over 3 million tons of whey. This requires over 25 000 m3 of installed membrane area and represents a large use of electrodialysis in the food industry [96]. Whey demineralizitation uses a conventional electrodialysis system, where cationand anion-exchange membranes are arranged in alternate order to form cell pairs. The most common feed for these systems is preconcentrated sweet whey (18–28%) [94, 95] In other commercial applications, acid whey, skim milk, reduced-lactose whey, milk and whey ultraﬁltration permeates are used as feed materials. Removing ash from whey with electrodialysis produces whey with up to 90% demineralization [8, 41, 98, 99]. The limiting factor in the demineralization of whey is the decrease in electrical conductivity of low ionic concentration solutions. For instance, the conductivity of whole milk is about 5 mS cm1, while fully

4.3 Electrodialysis Applications in the Food Industry

demineralized milk and whey have negligible electrical conductivity [64]. Concentrated whey as an electrodialysis feed is preferable because of its higher ionic concentration. When the ultraﬁltration permeate has been concentrated by a factor of four by reverse osmosis, the electrical conductivity increases by a factor of about two [100, 101]. Low conductivity is not wholly related to ionic concentration because the presence of lactose also depresses solution conductivity. Moreover, concentrated whey streams have a high protein concentration in the feed and this leads to a higher potential for protein-caused membrane fouling. A pH close to the protein isolectric point gives a better demineralization rate. The conversion of calcium salts to their ionized form by acidiﬁcation and deacidiﬁcation increases the conductivity of the solution. If calcium ions are replaced by sodium ions in the electrodialysis stack of deproteinated whey, the demineralization rate increases. The mobility of calcium ion is about 20% higher than the mobility of sodium ion [102, 103]; however, calcium ions have a tendency to form complexes with proteins and other species and these complexes tend to foul the membranes. The demineralization rate of whey in a good electrodialysis system is proportional to the conductivity to the power 0.95 [58, 102, 104]. Temperature also controls the demineralization of whey. For instance, batch-mode electrodialysis of ultraﬁltration permeate from casein whey to a 90% demineralization product requires different times at different feed temperatures. At 20 C, 90% demineralization takes 12 min, at 30 C it takes 8 min and at 40 C, only 6 min [8, 41, 104, 105]. In whey demineralization by electrodialysis, the ion removal rate follows ﬁrst-order kinetics for times up to 10–20 min. After that period, the curve of demineralization rate versus time ﬂattens. For whey demineralization by electrodialysis, the power requirement is 0.5 kWh lb1 of dried whey. The current efﬁciency of such systems is about 60–90% depending on the system-cleaning procedures [41, 99]. Demineralization of skim milk by electrodialysis reduces the level of ash and increases the calcium/phosphate ratio in skim milk powder. Electrodialysis demineralization of skim milk increases the stability of frozen skim milk and concentrated skim milk proteins [8, 36, 41, 99, 103]. For instance, the removal of about 40% of calcium ions by electrodialysis increases the shelf life of protein stored at 8 C from 1 to 17 weeks. A 70% calcium removal increases shelf life to 53 weeks [64, 100, 102, 104, 106–108]. Whey-protein concentrates are sometimes mixed with lactose (but not fat milk solids) to produce infant formula. It has been suggested that the commercial value of whey permeate can be increased by fermentation to lactic acid. The fermentation is carried out with a mixed culture of Lactobacillus helveticus and Streptococcus thermophilus [36, 38, 47, 99, 109–111]. This fermentation exhibits product inhibition; therefore, it is desirable to extract the lactic acid continuously as it is produced. Continuous lactic-acid production from whey permeates is done in a three-unit process comprised of the bioreactor, ultraﬁltration module, and electrodialysis cell. The ultraﬁltration module recycles all or part of the biomass back into the bioreactor and removes low molecular weight metabolites such as sodium lactate, which is a fermentation inhibitor. The sodium lactate solution is then extracted and concentrated continuously in an electrodialysis subsystem. The fermentation product without an electrodialysis subsystem yields an acid concentration of 40 g l1

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[52, 72, 90, 94, 112–115]. Adding an electrodialysis unit increases the ﬁnal lactate solution concentration to 130 g l1 [73, 94, 96, 116–118]. Electrodialysis after ultraﬁltration can extract 90% of the lactic acid from the fermentation bioreactor product. Sodium hydroxide is produced during the concentration of acid by electrodialysis [37, 43, 47, 94, 96]. However, continuous lactic-acid production has some potential disadvantages. Clogging of the ultraﬁltration subsystem membranes with protein deposits results in a drastic restriction of permeate ﬂow. In addition, the elimination of cationic ions from the fermentation broth changes the pH of the broth. For maximum bioreactor production, the fermentation is usually carried out at an optimum pH, which is usually signiﬁcantly higher than the pKa of the acid being formed. An example of fermentation and lactic acid production in the dairy industry using bipolar membrane electrodialysis is shown in Figure 4.7 (process drawing based Nordahl et al., 1998). A sterilized medium such as whey permeate is mixed with protein-hydrolysing enzymes and the resultant mixture is then pumped to a continuous fermenter containing a mixture of Lactobacillus helveticus and Streptococcus thermophilus. Lactic acid is then produced in the fermenter. The fermentation product is ultraﬁltered and the retentate contains the bacterial culture and nonhydrolysed whey protein. Dissolved ions pass through the membrane and concentrate the lactic acid in the permeate. An adjustment of pH with ammonium hydroxide is necessary to neutralize the acid produced. There are two different approaches to purify lactic acid from the permeate solution. The ﬁrst approach for purifying lactic acid from permeate solution was proposed by Norddahl et al. in 1998 [90]. Permeate from the ultraﬁltration process is passed

Figure 4.7 Lactic-acid production using bipolar membrane electrodialysis, adapted from Norddahl et al. [113].

4.3 Electrodialysis Applications in the Food Industry

Figure 4.8 Lactic-acid purification.

through an ion chelating ion-exchange resin to remove divalent ions. This prevents irreversible precipitation of calcium salts on the surface of the membranes in later electrodialysis processes. The eluant from the ion-exchange process then is concentrated with a two-step electrodialysis. Conventional electrodialysis is used in the ﬁrst step to concentrate the salt solution. Bipolar membrane electrodialysis is used in the second step to convert the salts into lactic acid, inorganic acids, and ammonium hydroxide. Lactic acid and ammonium hydroxide are recovered in two different streams using three-compartment bipolar electrodialysis as shown in Figure 4.8. The fermentation broth is pumped through a feed compartment composed of cationic and anionic membranes. The bipolar membranes adjacent to the cathode and anode generate OH and H þ groups, respectively. The lactate ions migrate toward the anode through the anion-exchange membrane and emerge into the product stream. Cations (such as ammonium, potassium, and sodium ions) migrate toward the cathode through the cation-exchange membrane, react with OH groups to form bases, and are removed from the stacks in the alkali stream. The overall recovery rate of lactic acid is about 85–90%, depending on the amount of sugar added to the fermenter [11, 47, 103, 119, 120]. Finally, the lactic acid is further puriﬁed or concentrated to the desired concentration using a falling-ﬁlm multistage vacuum evaporator or compression evaporator.

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The second approach proposed for purifying lactic acid via electrodialysis is shown by the Norddahl group in 2001 [113]. The fermentation liquid is ultraﬁltered and then acidiﬁed as shown in Figure 4.7. If the pH is above 3.8, it is usually lowered to 2.5–3.0, which is lower than the pKa value of lactic acid (3.86). The resultant free lactate ions bind with hydrogen ions to form lactic acid, having no net electrical charge. The low pH solution is then sent to nanoﬁltration or reverse osmosis to retain calcium and magnesium ions and molecules of molecular weight larger than 180. The calciumand magnesium-free permeate is then treated by three-compartment bipolar membrane electrodialysis. The bipolar membrane conﬁguration for lactate separation can also be a two-compartment conﬁguration where either cationic or anionic membranes are omitted; these possible conﬁgurations are shown in Figures 4.9a and b, respectively. In the two-compartment systems, only cations or anions are removed from the feed compartment and replaced with either protons or hydroxide ions. There is no concentrate stream in this mode. However, two-compartment bipolar membrane electrodialysis only partially deionizes the feed, since only cations or anions are removed. Bipolar membrane electrodialysis is simple and inexpensive as compared with other methods. Bipolar membrane electrodialysis systems boast lactic acid recovery rates as 90–98% based on the amount of sugar added to the fermenter [36, 43, 47, 96, 119]. The advantages of electrodialysis process in lactic acid production over conventional lactic acid production are: 1) 2) 3) 4) 5) 6)

no chemicals are needed to regenerate ion-exchange materials; the system has a higher operating efﬁciency; the system is easier to control; all of the efﬂuent or deplete streams are recycled; acids and bases are generated from optional bipolar membrane electrodialysis; the concentrate from nanoﬁltration, the only waste stream containing only calcium/magnesium ions and color compounds, is greatly reduced.

Bipolar membrane electrodialysis can also adjust the pH of dairy products. The dairy solution is circulated on the cationic side of the bipolar membrane where H þ ions are generated to lower the pH of the solution. Similarly, the solution is circulated on the anionic side of the bipolar membrane where OH ions are generated to increase the pH. The recommended current density is between 20–200 A m2, depending on the product to be treated [43, 59, 96, 98, 103, 115, 121–126]. This pH adjustment process simpliﬁes production technology, reduces cost, and eliminates the risk of explosion [59, 102, 104, 118, 121, 123, 127–133]. Recently, bipolar membrane electrodialysis has been applied to the puriﬁcation of dairy wastewaters using a three-stage process (Figure 4.10). In the ﬁrst stage, the wastewater is pretreated with a base to adjust the pH from 7 to 10. The base treatment partially precipitates the Ca2 þ , Mg2 þ , and PO43 ions that are present [36, 37, 67, 96, 108, 124]. In the second stage of the process, the pretreatment wastewater is then fed to a fermentation process where the lactose and other sugars present are converted to lactic acid using the bacteria Lactobacillus helveticus and Streptococcus thermophilus. A cell recycle stream circulates a stream from the fermenter through the microﬁltration unit and back into the fermenter. In the third stage, the permeate from the microﬁltration is fed to the electrodialysis

4.3 Electrodialysis Applications in the Food Industry

Figure 4.9 Lactate purification (a) two- and (b) three-compartment configuration.

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Figure 4.10 Purification of dairy wastewater, adapted from Boergardts et al. [114] (data in Boergardts et al.).

system through either a nanoﬁltration unit or a selective ion exchanger to remove any residual ions. In the ﬁnal stage, the concentration of lactic acid in the wastewater is reduced. The produced wastewater exhibits a low chemical oxygen demand (COD) load. Bipolar membrane electrodialysis allows the isolation of free acid in high concentration, and by-product alkali is utilized to elevate the pH in the pretreatment stage. The bipolar membrane electrodialysis process in wastewater treatment is shown in Figure 4.10. The fermentation broth passes through the diluate compartment. The pH in each of the two end compartments is controlled by a bipolar membrane. The lactic acid is removed from the feed stream by the anion-exchange membrane, while the alkali ions are removed through the cation-exchange membrane. In order to achieve both high lactic acid concentration and low COD concentration [55, 114], Boergardts et al. suggest that concentration of fermentation broth is needed to lower wastewater concentration. During the electrodialysis step, water is circulated through the membrane stack and the products (lactic acid and alkali) are removed from the system continuously at constant concentration. Boergardts et al. [114] also suggest that the electrodialysis process can be carried out by a two-stage process. The fermentation broth is pumped through the ﬁrst stage of bipolar membrane electrodialysis, where the broth is continuously diluted to an ionic concentration of about 10–15 g l1. In the second stage, a conventional electrodialysis system is used to further reduce the concentration of ions in the wastewater stream. The sodium lactate from the second electrodialysis stage can be recycled to the ﬁrst electrodialysis stage to increase the feed concentration of lactate as shown in Figure 4.10. Results show that the COD concentration is reduced by 85–95%, the free lactic concentration is about 200 g/l, and the alkali solution concentration is about 2 mol l1 [11, 35, 36, 44, 47, 98, 99, 119]. 4.3.4 Protein Fractions

The protein fraction recovery process is similar to whey demineralization and is based on the two primary characteristics of electrodialysis: decreasing the ionic

4.3 Electrodialysis Applications in the Food Industry

concentration (desalting) and increasing the ionic concentration (salting-out effect). These two characteristics are used in numerous applications to remove impurities that are insoluble in high or low ionic strength or in the selective removal of proteins of interest. One the ﬁrst protein fraction technologies was developed in 1982 for the separation of enriched b-lactoglobulin (b-lg) and a-lactalbumin (a-la) fractions from whey [72]. Ultraﬁltration is used to concentrate the whey proteins and to partially remove water, salts, lactose, and other low molecular weight compounds. The permeate from ultraﬁltration is adjusted to a pH of 4.65 with either HCl or NaOH [95]. The following electrodialysis demineralization step removes low molecular ions such as sodium, potassium, calcium, and magnesium. The demineralized concentrate is readjusted to a pH of 4.65 if necessary either with 0.1 N HCl or NaOH. b-lg precipitates in this step [37, 94, 96] and the precipitate is separated from the solution by centrifugation. Using this method, the protein solutions are desalted with minimal loss of solute. About 33% of the acid whey proteins are recovered by using pH adjustment coupled with electrodialysis. In 1995, Stack et al. [134] developed a new process using thermal treatment coupled with the previously described protein separation methods. Stacks process was based on an earlier process developed by Pearce et al. [121], a well-known process based on the thermal separation of whey proteins. In the Pearce process, the raw material is treated to reduce its speciﬁc gravity and ionic strength to levels less than 25% of the original values. Next, a-la is aggregated for 30 s by heating the whey to 55–70 C. The ﬂocculated a-la is recovered by centrifugation, whereas the soluble b-lg remains in the whey solution with other constituents. Stack et al. extended this concept to develop an efﬁcient integrated process for treating whey and recover its constituents, especially pure b-lg fraction, the enriched a-la fraction, and lactose, as shown in Figure 4.11. In the ﬁrst step, the whey is treated to reduce its mineral content using electrodialysis to achieve 70% demineralization. The cation exchange resin column removes the rest of potassium, sodium, magnesium, and particularly calcium. The treated whey is then subjected to a heat treatment at between 71 and 98 C for 50–95 s.

Figure 4.11 Recovery of proteins from whey, adapted from Stack et al. [134] (data in Stack et al.).

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At this temperature and these ionic conditions, the b-lg remains soluble in the solution. After the heat treatment, the proteins in whey are rather soluble. The whey is then concentrated by a two-stage process to between 55–63% and the lactose crystallizes as the concentrated solution cools. In the second stage, the pH of the whey solution is adjusted to between 4.3 and 4.7 at a temperature less than 10 C and is then heated to 35–54 C for 1–3 h. The a-la component of the solution ﬂocculates. Stack et al. [134] did not report the yield of either b-lg or a-lac proteins in whey fractions. Combined with cation-exchange membranes, bipolar membrane electrodialysis can lower the pH of the solution in the compartment next to cationic side of the bipolar membrane. Bazinet and coworkers [41, 47, 98, 135] also fractionated whey proteins with bipolar membrane electrodialysis. As the whey solution circulates through the cells, the pH of the solution is lowered from 6.9 to 4.6. A Feed of 5% protein concentration, processed with the bipolar membrane electrodialysis system, produced a 98% pure b-lg fraction with a 44% recovery. A feed of 10% protein concentration is optimum for the bipolar membrane electrodialysis system, and a 95.3% pure b-lg fraction at 53.4% recovery can be achieved at that feed concentration. The b-lg-enriched fraction contains 2.7% of a-la for 98% total protein purity [120]. The performance of bipolar membrane electrodialysis is improved as the initial concentration of protein increases. However, if the initial concentration is above 10%, the conductivity of the solution becomes a limiting factor. Conventional electrodialysis and bipolar membrane electrodialysis show advantages in protein fractionation compared to conventional heat-treatment methods. The electrodialysis systems give rapid and controlled recovery of salts without diluting the product. The very low molecular weight protein and peptides can be easily demineralized. The electrodialysis processes for protein fractionation are well suited for recycling the salts responsible for the salting-out effect. Both electrodialysis and bipolar membrane electrodialysis can concentrate salts in one stream, while desalting the other stream.

4.4 Hybrid Technologies 4.4.1 Electrodeionization

Based on the concepts of both electrodialysis and ion exchange resin columns, electrodeionization is the membrane process in which cation and anion exchange resins are packed between the two membranes in the feed compartment to enhance the transport of ions across the system (Figure 4.12). Electrodeionization has been widely used for ultrapure water production because it requires less energy than electrodialysis systems at low ionic concentration [9, 52]. However, electrodeionization has disadvantages that must be overcome. Since the system contains ionexchange resins rather than a spacer between the two membranes, system leakage

4.4 Hybrid Technologies

Figure 4.12 Electrodeionization.

can be severe, which greatly reduces system performance. Moreover, electrodeionization systems often exhibit uneven ﬂow distribution due to ﬂow channels created by the resins packed between the two membranes. These two disadvantages drive the cost for high-performance electrodeionization systems to the point that application of the technology is currently limited to ultrapure water production. Arora et al. [136] developed a method to bind the resins together to form a wafer that was suitable for the recovery of lactic acid. The wafer contains not only the properties of the ion-exchange resins, but also the function of the spacer; therefore, wafer-enhancedelectrodeionization technology has the potential to lower system costs. Moreover, with lower power costs for the separation of ions at low concentrations, waferenhanced-electrodeionization could separate low conductivity solutions found in food processing, such as milk and juice. 4.4.2 Electrochemical Coagulation

Water electrolysis, the formation of a boundary layer at the electrode/solution interface, and a convection-diffusion phenomenon are basic concepts for electrochemical coagulation (Figure 4.13) [11, 35, 135]. The pH increase in the anode compartment and decrease in the catholic compartment are results of decreased ion transport across the membrane. When a membrane separates the compartments, there is an increase in the acidity or alkalinity with respect to the bulk solution while, without the membrane, the increase in acidity and alkalinity only happens at the boundary layers formed at the electrode/solution interfaces. Because acid and base are created at the anode and cathode, respectively, the rinse solutions in these two compartments can be used for juice and dairy treatment. For instance, the low-pH solution generated from the anode can be used as the treatment solution for precipitating whey protein, especially a-lactalbumin. Another application is to use the cathode rinse to clean the membranes and the electrodialysis stack.

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Figure 4.13 Electrochemical coagulation.

The high-pH solution from the cathode chamber can be used to balance the acidity of juices such as pineapple, orange, or grape [135]. 4.4.3 Electroreduction

The use of electric voltage to break covalent bonds, thereby forming new molecules, is the basic concept of electroreduction technology. The covalent bond is broken by the electrical ﬁeld, while the solution is circulating in the cathode compartment of the electrodialysis stack [11, 35, 94]. A new bond is formed when the solution moves out of the compartment, as shown in Figure 4.14. The breaking of divalent bonds, especially disulﬁde bonds in proteins, has been applied widely in protein analysis of biological species. This same phenomenon could be applied to protein separations in the dairy industry. By using electric potential, the disulﬁde bonds of a-la and b-lg can be broken, and new chains of proteins with different side chains can be formed so that proteins of higher purity can be recovered. Using electroreduction, the production of free sulfhydryl (SH) groups and prevention of the thiol-disulﬁde interchange reaction increases the stability of proteins.

4.5 Conclusion and Future Innovations

Many applications of electrodialysis are found throughout the food industry. The electrodialysis techniques used include conventional electrodialysis,

4.5 Conclusion and Future Innovations

Figure 4.14 Electroreduction.

three-compartment electrodialysis, bipolar membrane electrodialysis, and other hybrid technologies. However, the commercial use of electrodialysis techniques is still limited to niche applications such as juice deacidiﬁcation or whey-protein demineralization. This nonacceptance is attributed to the fact that the mechanisms of electrolytic phenomena are very complex, especially for multicomponent systems. The lack of a detailed understanding of oxido-reduction in electrolytic phenomena and redox reactions of food compounds limits the broad application of electrodialysis and the electrolytic cell. Electrodialysis and electrolytic cell techniques have the potential to improve and integrate into more food processes. Possible candidate applications include the selective removal of ions, waste recovery, and others awaiting exploration. Although electrodialysis has matured during the past several decades, its application in the food industry is still limited. The food industries are typically late adopters of new technology as compared to other industries such as the chemical or pharmaceutical industries. Some factors that have prevented electrodialysis from wide acceptance in the food industries are membrane fouling, limited cleanability, and poor membrane chemistry. For food applications, membrane fouling is a severe problem that decreases the system performance and increases the cost per amount of product. As addressed earlier, electrodialysis membranes are typically cleaned by the system vendor rather than though in-situ cleaning. This makes the technology inconvenient and costly. In electrodialysis, ion removal is usually nonselective, which limits the applicability of electrodialysis for speciﬁc ion removal. These factors provide many opportunities for electrodialysis research. Innovation of new membrane chemistry designed for low fouling and high selectivity in ionic removal would expand the use of electrodialysis. Moreover, innovative system design leading to easier installation and in-situ cleaning is a key requirement for the expansion of the use of electrodialysis for food–related applications. For more information on electrodialysis purchase, a vendor list is given in Appendix 4.A.

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Appendix 4.A: Electrodialysis Vendor List

Company name

Location

Contact information

Alpine Technical Services, Inc. Ameridia Applied Membranes, Inc. Applied Water Solutions, Inc.

Utah, USA

www.alpinetech.US

New Jersey, USA California, USA Massachusetts, USA

Baymont Technologies, Inc. CelTech, Inc. ChemTreat, Inc. Crane Environmental Eden Puriﬁcation Systems

Texas, USA North Carolina, USA Virginia, USA Pennsylvania, USA Connecticut, USA

Eurodia Industrie Exergy Tecnologies Corp. GE Water and Process Technologies Ion Power, Inc. Jinan Haochua Industry Co., Ltd. Koch Membrane Systems, Inc Minntech Corporation Sparkling Clear Industries TTS Technologies

New Jersey, USA California, USA Pennsylvnia, USA

www.ameridia.com 760-727-3711 www.appliedwatersolutions. com 281-260-0667 www.celtechinc.com www.chemtreat.com 732-202-9211 www.edenpuriﬁcationsystem. com www.eurodia.com 949-679-3990 www.gewater.com

Delaware, USA Shandong, China

www.ion-power.com www.jnhaohua.com

Massachusetts, USA

978-657-5208

Minnesota, USA Texas, USA Tampere, Finland

www.mintech.com www.sparklingclear.com 358-3-31422011

Nomenclature

C c cþ E F I k Q t tþ u v V Z

Ion concentration Concentration of anions Concentration of cations Energy consumption Faraday constant Electric current Equilibrium constant Flow rates Transport number for anions Transport number for cations Velocity of cations Velocity of anions Voltage potential applied across the stack Valence of ions

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5 Membrane Processes in Must and Wine Industries Maria Norberta De Pinho In wine industries the conventional processes of ﬁltration for clariﬁcation and cold treatment for tartaric stabilization are giving place to an increasing use of alternative membrane processes namely micro/ultraﬁltration (MF/UF) for clariﬁcation and electrodialysis (ED) for tartaric stabilization. This wide use of membrane processes is carried out, most of the times, having in mind a single operation of application. For example, the clariﬁcation by MF or UF is optimized in terms of productivity and of preservation of the organoleptic properties like ﬂavors and aromas. However, the removal of macromolecules like polysaccharides and polyphenols not only has a crucial importance on the organoleptic properties but also plays an important role on the wine tartaric stability and therefore in the subsequent operation of Electrodialysis (ED). For that reason, the integration of these operations will be the object of analysis. Nanoﬁltration (NF) is assessed as a fractionation technique for the simultaneous concentration and rectiﬁcation of grape musts.

5.1 Introduction

Membrane operations are nowadays an essential part of the wine-making process. As shown in Figure 5.1 after must fermentation the clariﬁcation operation is associated with tangential microﬁltration (MF)/ultraﬁltration (UF) and the tartaric stabilization operation is associated to electrodialysis (ED). Wine clariﬁcation, traditionally carried out by diatomaceous-earth ﬁltration, is being replaced by tangential microﬁltration (MF) and ultraﬁltration (UF). Besides their advantages on the continuous and automatic mode of operation they brought enormous environmental beneﬁts on the elimination of solid wastes of diatomaceousearth ﬁltration media and microorganisms. In wine tartaric stabilization the complex conventional sequence of wine cooling, tartrate crystal seeding, dynamic crystallization and diatomaceous-earth ﬁltration is being replaced by electrodialysis (ED) [1–3]. In parallelwiththeoperatingadvantagesofbeinganeasyandcontrollableprocess thereare beneﬁts of energy savings and of no generation of large amounts of diatomaceous-earth solid wastes that constitutes an important asset from the environmental point of view.

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Production flow-sheet

Complementar Operations

Grapes

• Addition of Concentrated Must (CM) or Rectificated Concentrated Must (RCM)

Crushing Pressing Grape Must

Enological Practices • Chaptalization - only allowed in Northern Europe

• Partial Concentration in the Winery

Enrichment Rectification and Acid Correction

Fermentation

• Addition of Acid Tartaric or Electrodialysis with Bipolar Membranes • Sedimentation and Racking • Protein fining • Centrifugation • Kieselguhr or Diatomaceous-Earth filtration • Dead end Microfiltration

Wine

Botling

Clarification

• Tangential Micro-and Ultrafiltration

Tartaric Stabilization

• Cold Crystalization process with or without crystal seeding

Biological Stabilization

• Electrodialysis

• Ion Exchangers of styrene and vinyl benzene

• Pasteurization • Use of Sulfur Dioxide • Sterile Filtration

Figure 5.1 Grape must and wine production.

One of the facts behind the restricted application of MF/UF to wine clariﬁcation has to do with the lack of knowledge on the possible removal of polysaccharides or other macromolecules that may be of major relevance to the wine quality. At the same time, these macromolecules play an important role on the tartaric stability. Moreover, the colloids removed by clariﬁcation may act as natural inhibitors of potassium hydrogen tartrate precipitation. The integration of MF/UF with ED becomes therefore of crucial importance and will be the object of concern in the analysis that follows [4, 5]. As shown in Figure 5.1, the grape must, prior to the fermentation, may be subjected to enrichment and acid correction by addition of rectiﬁed concentrated must and tartaric acid, respectively. The substitution of these operations by membrane operations like nanoﬁltration and reverse osmosis is the object of research [6– 8] and again one should view these operations further integrated with those of wine processing and namely with electrodialysis for tartaric stabilization.

5.2 Wine Clarification by Microfiltration and Ultrafiltration

Feuillat [9] claims that wine turbidity is caused by suspended material like yeast residues and macromolecular compounds with colloidal behavior. The clariﬁcation operation, performed to remove these compounds, is assessed both in terms of

5.2 Wine Clarification by Microfiltration and Ultrafiltration

productivity and polysaccharide removal. Serrano et al. [10] compared traditional ﬁltration with tangential MF and concluded that tangential microﬁltration led to wines with lower polysaccharide content. The membrane fouling, besides having direct consequences on MF and UF productivity, brings additional problems related to the removal of macromolecules essential to wine quality. Belleville et al. [11, 12] identiﬁed some polysaccharide as major responsible for the fouling of MF mineral membranes. Cameira-dos-Santos [13] and Vernhet et al. [14, 15] have proved that polysaccharides and polyphenols also play an important role in the fouling of MF organic membranes. In the prespective of optimizing productivity and minimizing polysaccharide removal, tangential microﬁltration and ultraﬁltration are analysed in Figures 5.2 and 5.3 as a function of the operating parameters of transmembrane pressure and concentration factors. The membranes used are: a MF membrane with 1 mm pore size and a UF membrane with the molecular weight cut-off (MWCO) of 100 kDa. Both membranes are made of a ﬂuorpolymer and are supplied by Alfa Laval – Denmark (former DSS-Denmark). Figure 5.2 displays the productivity or the permeate ﬂuxes decline in microﬁltration of a white wine, Vinho Verde (Portugal) [4, 16], versus the concentration factor. An increase of the transmembrane pressure from 0.6 to 1.0 bar means a signiﬁcant gain in the permeate ﬂuxes. Figure 5.3 displays the productivity or the permeate ﬂuxes decline in ultraﬁltration of a white wine, Vinho Verde (Portugal) [4, 16], versus the concentration factor. The permeate ﬂuxes are practically independent of the transmembrane pressure and no gain in productivity is obtained for transmembrane pressures higher than 1.0 bar.

Figure 5.2 Variation of white wine microfiltration permeate fluxes, Jv (L h1 m2), with the concentration factor. Transmembrane pressures ranging from 0.6 105 to

1.4 105 Pa. MF membrane – FSM1.0PP – with 1.0 mm pore size. Experiments run in a plate and frame DDS Lab-Unit, type 20, with 0.036 m2 of membrane surface area.

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Figure 5.3 Variation of white-wine ultrafiltration permeate fluxes, Jv (L h1 m2), with the concentration factor. Transmembrane pressures ranging from 1.0 105 Pa to

2.6 105 Pa. UF membrane – FS40PP – with MWCO of 100 kDa. Experiments run in a plate and frame DDS Lab-Unit, type 20, with 0.036 m2 of membrane surface area.

Figure 5.4 displays the productivity or the permeate ﬂuxes decline in microﬁltration of a red wine, Vinho Verde (Portugal) [4, 16], versus the concentration factor. The permeate ﬂuxes are practically independent of the transmembrane pressure. Figure 5.5 displays the productivity or the permeate ﬂuxes decline in ultraﬁltration of a red wine, Vinho Verde (Portugal) [4, 16], versus the concentration factor. The

Figure 5.4 Variation of red-wine microfiltration permeate fluxes, Jv (L h1 m2), with the concentration factor. Transmembrane pressures ranging from 0.6 105 to

1.4 105 Pa. MF membrane – FSM1.0PP – with 1.0 mm pore size. Experiments run in a plate and frame DDS Lab-Unit, type 20, with 0.036 m2 of membrane surface area.

5.2 Wine Clarification by Microfiltration and Ultrafiltration

Figure 5.5 Variation of red-wine ultrafiltration permeate fluxes, Jv (L h1 m2), with the concentration factor. Transmembrane pressures ranging from 1.0 105 to

2.6 105 Pa. UF membrane – FS40PP – with MWCO of 100 kDa. Experiments run in a plate and frame DDS Lab-Unit, type 20, with 0.036 m2 of membrane surface area.

permeate ﬂuxes are dependent on the transmembrane pressure and its increase leads to a signiﬁcant productivity gain. The permeate ﬂuxes of MF and UF of a red wine are much lower than those of MF and UF of a white wine. At the transmembrane pressure of 1.0 bar, the MFand the UF of a white wine yields ﬁnal permeate ﬂuxes of 118 and 129 L h1 m2, respectively. At the same transmembrane pressure of 1.0 bar, the MF and the UF of a red wine yields ﬁnal permeate ﬂuxes of 34 and 18 L h1 m2, respectively. For white and red wine, the removal of polysaccharides and polyphenols in the operations of MF and UF is shown in Tables 5.1 and 5.2, respectively. The wine clariﬁcation by microﬁltration is associated with a small removal of polysaccharides and polyphenols for the case of white wine and to a slightly higher removal for the case of red wine. Table 5.1 Clarification by microfiltration.

Wine

DP (bar)

Percentage of removal Polysaccharides (%)

Polyphenols (%)

White

0.6 1.0 1.4

11.4 7.6 7.7

2.1 0.9 2.6

Red

0.6 1.0 1.4

24.6 22.8 23.1

9.6 12.6 10.2

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Table 5.2 Clarification by ultrafiltration.

DP (bar)

Wine

Percentage of removal Polysaccharides (%)

Polyphenols (%)

White

1.0 1.8 2.6

16.4 16.4 18.7

0.0 0.8 4.0

Red

1.0 1.8 2.6

82.9 83.9 94.5

31.5 43.4 54.1

The clariﬁcation of white wine by ultraﬁltration also leads as in the case of MF to a low removal rate of polysaccharides and negligible removal of polyphenols. In contrast, for the case of red wine there is a signiﬁcant removal of polysaccharides and polyphenols. Upon the degree of fouling, the regeneration of MF and UF membranes is made through the circulation of water or solutions of detergent at different temperatures and circulation times. A cleaning sequence is composed of the following steps: 1) circulation of water at the temperature of 20 C and for 30 min; 2) circulation of water at the temperature of 50 C and for 30 min; 3) circulation of water at the temperature of 50 C and for 60 min; 4) circulation of Ultrasil11 solution at the temperature of 50 C and for 30 min; 5) circulation of Ultrasil11 solution at the temperature of 50 C and for 60 min; 6) circulation of Ultrasil11 solution at the temperature of 50 C and for 3 h. The different cleaning sequences yield the results shown in Tables 5.3 and 5.4 for microﬁltration and ultraﬁltration, respectively. They show along the different steps the percentage of permeate ﬂuxes recovery.

Table 5.3 Sequence of membrane regeneration operations after clarification by microfiltration.

Clarification

Cleaning procedure

Wine

DP (bar)

First step: water 20 C, 30 min

Second step: water 50 C, 30 min

White

0.6 1.0 1.4

83% 82% 80%

93% 104% 94%

— — —

Red

0.6 1.0 1.4

81% 48% 31%

73% 75% 57%

97% 97% 98%

Third step: Ultrasil11 0.5%, 50 C, 30 min

5.3 Wine Tartaric Stabilization by Electrodialysis ? Table 5.4 Sequence of membrane regeneration operations after clarification by ultrafiltration.

Clarification

Cleaning procedure

Wine

DP (bar)

First step: water 20 C, 30 min

White

1.0 1.8 2.6

93% 95% 92%

— — —

— — —

— — —

Red

1.0 1.8 2.6

27% 18% 15%

38% — —

75% 65% 44%

95% 91% 66%

Second step: water 50 C, 60 min

Third step: Ultrasil11 1%, 50 C, 60 min

Fourth step: Ultrasil11 1%, 50 C, 3 h

The easier process of regeneration is relative to ultraﬁltration of white wine and consists just on a single step of circulation of water at the temperature of 20 C and for 30 min. The microﬁltration of white wine requires a further step of circulation of water at the temperature of 50 C for 30 min. For red wine the regeneration process is more difﬁcult and in the case of MF it requires an additional step of cleaning through circulation of a solution with 0.5% of Ultrasil11 at the temperature of 50 C for 30 min. The UF of red wine leads to severe membrane fouling and to the need for circulating 1% Ultrasil11 solutions for longer times of 3 h. Moreover, if the UF operating pressures are as high as 2.6 bar, the membrane fouling is irreversible and the permeate ﬂuxes are only recovered to 66%.

5.3 Wine Tartaric Stabilization by Electrodialysis [4, 5]

Potassium hydrogen tartrate (KHT) is a natural constituent of grapes. Alcoholic fermentation during winemaking leads to a decrease in the KHTsalt solubility due to the presence of ethanol. As a consequence, at normal storage temperatures an untreated wine is supersaturated in KHT and undesirable precipitation can occur in the bottles. To overcome this problem, the cold tartaric stabilization method is traditionally used. As shown in Figure 5.6 this consists of a complex sequence of wine cooling, tartrate crystal seeding, dynamic crystallization and diatomaceous-earth

Figure 5.6 Cold tartaric stabilization process.

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electrode washing solution

Feed Brine + + + + +

Anode

+ + + + + + + + +

+ + + + + + + + + + + + corrente eléctrica electric current + + -

+ + + + + + + + + + + + + +

-

Cathode

Brine Diluate

Figure 5.7 Schematic representation of electrodialysis.

ﬁltration. Besides not allowing a precise control of the ﬁnal KHT concentration this method may lead to unwanted precipitation of polysaccharides and polyphenols together with KHT crystals. These limitations are overcome in the treatment by electrodialysis (ED), which is a method based on ion electrical migration in a singlestage operation, as shown in Figure 5.7. In ED the wine circulates in rectangular channels conﬁned by cation- and anionselective membranes and by the action of an external electric ﬁeld normal to the membranes, the ions are forced to migrate to the electrodes, giving rise to a wine stream depleted in ions [17]. This is schematically shown in Figure 5.7 where the wine circulates in the diluate compartments that alternate with the brine compartments. An important feature of ED is the fact that during wine circulation there is a reduced surface area of contact with the membrane walls of the diluate compartment. This is in contrast with the cold tartaric stabilization process that involves a ﬁltration step where the wine percolates through porous media of extensive surface areas and leads very often to adsorption of organic molecules of great relevance for the organoleptic properties of the wines. Also, the ED dense polymeric membranes are not prone to adsorption phenomena. The nonalteration of the organoleptic properties of the wines constitutes therefore a strong asset of ED. Another asset is the ﬂexibility in reaching any degree of KHT removal through the variation of the ED operating time.

5.4 Influence of MF/UF Polysaccharide Removal on Wine Tartaric Stability

Figure 5.8 Influence of electrodialysis operating time on the wine conductivity.

Figure 5.8 displays the variation of the wine conductivity with ED operating time. The decrease of conductivity is associated with the removal of potassium and tartaric acid as the cations and anions present in higher concentrations. The deionization degree (DEID) is deﬁned as: DEID ¼ ((initial conductivity–ﬁnal conductivity)/initial conductivity). In Figure 5.8, the DEID values are assigned in percentages at the points of sample collecting, full squares and full triangles, for white and red wine, respectively. At the various degrees of KHT removal, the wine tartaric stability is assessed through the determination of the saturation temperature, Tsat [5]. Figure 5.9 displays for a white wine the variation of the saturation temperature as a function of the degree of ED deionization. The experimental results are correlated by the equation: Tsat ¼ 20:30:44 deionization percentage

ð5:1Þ

5.4 Influence of MF/UF Polysaccharide Removal on Wine Tartaric Stability

After wine clariﬁcation by microﬁltration with a membrane of 1 mm pore size and ultraﬁltration with a membrane of 100-kDa MWCO, the permeate and concentrate streams were subjected to a polysaccharide precipitation process [4, 16]. The results

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Figure 5.9 Variation of the saturation temperature with the degree of ED deionization.

obtained together with the corresponding values for raw wine are presented in Table 5.5. The 10% polysaccharide removal during MF with a membrane of 1 mm pore size is relatively low when compared with 50.3% obtained by Serrano et al. [10] with a 0.4-mm organic membrane. A 16% polysaccharide removal is obtained with the UF membrane of 100 kDa. Escudier et al. [18] reported a value of 92% with a 20 kDa membrane and that led to a very unstable wine. The role of the polysaccharides on wine stability is assessed through the measurement of the crystallization induction times of potassium hydrogen tartrate on a model solution of ethanol, potassium hydrogen tartrate and tartaric acid in the same concentration as in raw wine and three model solutions prepared from the model solution and adding raw wine polysaccharides, UF permeate polysaccharides and UF concentrate polysaccharides. The crystallization induction times are determined by monitoring the conductivity of a solution while lowering the temperature to a pre-set value, in order to induce salt precipitation. After an initial decay, the conductivity stabilizes in a plateau and then decreases again when precipitation starts. The time interval between this instant and the instant when temperature reaches the pre-set value is the induction time. The results are displayed in Table 5.6.

Table 5.5 Variation of polysaccharides content in raw wine and wine clarified by MF and UF.

Polysaccharides (mg l1)

Raw wine Permeate Concentrate

MF

UF

334 300 665

334 281 800

5.5 Nanofiltration of Grape Must for Sugar/Organic Acids Fractionation Table 5.6 Influence of polysaccharides of UF streams on KHT crystallization induction time.

Polysaccharides (mg L1)

Induction time (h)

0 30.2 30.0 30.8

14.3 20.3 22.0 35.6

Model solution Model solution with raw wine polysaccharides Model solution with UF permeate polysaccharides Model solution with UF concentrate polysaccharides

The induction times obtained with UF permeate polysaccharides are slightly higher than those obtained with raw wine polysaccharides. The UF concentrate polysaccharides led to higher induction times and therefore showed a higher inhibition effect.

5.5 Nanofiltration of Grape Must for Sugar/Organic Acids Fractionation

Grape must quality is of major importance in the deﬁnition of the wine character. Enrichment of must prior to fermentation is one process that is used to overcome reduced levels of sugars in a particular vintage. As shown in Figure 5.1 this is traditionally done by adding sucrose from beet and cane sugar or grape musts – concentrated must (CM) and rectiﬁed concentrated must (RCM). The vacuum evaporation (VE) is used to produce CM and is very often associated to the depletion of varietal aromas and to the production of off-ﬂavors [19, 20]. More recently, reverse osmosis (RO) is being used for must concentration [6, 21]. However, if must rectiﬁcation is considered, an additional operation of ion exchange is required and that brings severe ecological problems due to the need for resin regeneration and its disposal [22]. Rosa Santos et al. [8] propose nanoﬁltration for the simultaneous concentration and rectiﬁcation of grape must. This is investigated through the capability of NF to fractionate sugars from the organic acids in a grape must from Table 5.7 Composition of the grape must model solutions and the EDM grape must.

Grape must model solutions RD1

Tartaric acid (g L1) Malic acid (g L1) Total sugar (g L1) pH Conductivity (mS cm1)

RD2

EDM

RD1T

RD1M

RD2T

RD2M

2.0 — 150 2.64 892

— 2.5 150 2.41 700

2.6 — 200 2.52 913

— 3.3 200 2.33 685

2.0 5.0 107 3.19 2200

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the region of Vinho Verde production (Entre Douro e Minho (EDM), Portugal) and four model solutions of grape must. The composition of the grape must designated by EDM and of the four model solutions is shown in Table 5.7. The grape must model solutions were prepared as described by Rosa Santos et al. [8] and designated by RD1T, RD1M, RD2T and RD2M. The nanoﬁltration is performed with a NF 270 membrane supplied from FilmTec (Minneapolis, MN) and yields the results displayed in Figure 5.10. For the model solutions, the gap between the rejection coefﬁcients to the sugars –glucose and fructose – and to the acids – tartaric and malic acids – is very pronounced. The sugars being rejected more than 88% and the acids less than 37% means that the major part of the sugars are retained in the NF concentrate stream and the organic acids permeate preferentially to the permeate stream. This demonstrates the NF capability for sugars/organic acids fractionation in grape musts. This fractionation is enhanced with the increase of the total sugar content from 150 g L1 in the RD1T and RD1M to 200 g L1 in the RD2T and RD2M.

NF270 - RD1T

NF270 - RD1M

100 80

100 90.2

92.7

80

60

92.5

60

40

40

20

37.3

20

30.8

0

0 Glu

Fru

TA

Glu

MA

Fru

TA

MA

NF270 - RD2M

NF270 - RD2T

100

100 80

94.7

88.2

80

90.9

60

60

40

40

20

88.7

91.7

20

22.9

0

25.8

0 Glu

Fru

TA

MA

Glu

Fru

TA

MA

NF270 - EDM

100

Glu – glucose

80

Fru – Fructose

60

TA – Tartaric Acid

98.0

98.3

89.1 64.3

40 20

MA – Malic Acid

0 Glu

Fru

TA

MA

Figure 5.10 NF Rejection coefficients to glucose, fructose, tartaric acid and malic acid in grape musts. Membrane NF270.

References

For the EDM grape must, the gap between the rejection coefﬁcients to the sugars and to the acids is less pronounced. Among the acids, there is a preferential permeation of malic acid.

Acknowledgments

Thanks go to Dr Rui Viegas for helpful discussions and to the ﬁnancial support of the Foundation for Science and Technology (FCT) through POCI/AGR/58450/2004 and REEQ/764/EQU/2005.

References 1 Escudier, J.-L., Saint Pierre, B., Batlle, J.-L.,

2

3

4

5

6

and Moutounet, M. (1993) Procedé et dispositif automatique de stabilisation tartrique de vins. French Patent FR 2709308-A1 Gonçalves, F., Cameira dos Santos, P.J., Spranger, M.I., Pereira, O.M., Santos, F., and Pires da Silva, M. (1998) Essais de stabilisation tartrique sur Vinho Verde: Étude comparatif de l électrodialyse et dun procédé traditionnel. in Proceedings of XXIII eme Congrés Mondial de la Vigne et du Vin, Lisbon, Portugal. Gonçalves, F., Cameira dos Santos, P.J., Barros, P., Pinheiro, J.A., and Pinho, M.N. (1999) Application de lélectrodialyse a la stabilisation tartrique de vin de Porto. Comparaison avec le procedé de froid par stabulation longue. in Proceedings of 6e Symposium International dOEnologie, Bordeaux, France. Gonçalves, F. (2002) Optimizaç~ ao da clariﬁcaç~ao e Estabilizaç~ ao Tart arica de Vinhos por Processos de Membranas. Inﬂu^encia das Macromoléculas e da Composiç~ao Iónica, Instituto Superior Técnico, Universidade Técnica de Lisboa, Portugal. Gonçalves, F., Cameira dos Santos, P.J., Barros, P., Pinheiro, J.A., and Pinho, M.N. (2003) Wine tartaric stabilization by electrodialysis and its assessment by the saturation temperature. J. Food Eng., 59, 229–235. Mietton-Peuchot, M., Milisic, V., and Noilet, P. (2002) Grape must concentration by using reverse osmosis.

7

8

9

10

11

12

13

Comparison with chaptalization. Desalination, 148, 125–129. Pinho, M.N., Geraldes, V., and Catarino, I. (2007) Method for Simultaneous Concentration and Rectiﬁcation of Grape Must using Nanoﬁltration and Electrodialysis, PCT PT 2007/000045. Rosa Santos, F., Catarino, I., Geraldes, V., and de Pinho, M.N. (2008) Concentration and rectiﬁcation of grape must by nanoﬁltration. Am. J. Enol. Viticult., 59 (4), 446–450. Feuillat, M. (1987) Stabilisation et clariﬁcation des vins: aspects colloıdaux. Rev. Oenol., 45, 8. Serrano, M., Pontens, B., and RibéreauGayon, P. (1992) Étude de différents membranes de microﬁltration tangentielle. Comparaison avec la ﬁltration sur précouche de diatomées. Journal international des sciences de la vigne et du vin, 26, 97. Bellevillle, M.-P., Brillouet, J.-M., Tarodo de la Fuente, B., and Moutounet, M. (1990) Polysaccharide effects on cross-ﬂow microﬁltration of two red wines with a microporous alumina membrane. J. Food Sci., 55, 1598. Bellevillle, M.-P., Brillouet, J.-M., Tarodo de la Fuente, B., and Moutounet, M. (1992) Fouling colloids during microporous alumina membrane ﬁltration of wine. J. Food Sci., 57, 396. Cameira-dos-Santos, P. (1995) Colmatage en microﬁltration tangentielle: mise en évidence dinteractions entre les polysaccharides et les polyphénols du vin

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14

15

16

17

et des membranes polymeriques, PhD Thesis, École Nationale Supérieure Agronomique de Montpellier, France. Vernhet, A., Bellon-Fontaine, M.-N., Brillouet, J.-M., Roesink, E., and Moutounet, M. (1997) Wetting properties of microﬁltration membrane: determination by means of the capillary rise technique and incidence on the adsorption of wine polysaccharide and tannins. J. Membrane Sci., 128, 163. Vernhet, A., Pellerin, P., Belleville, M.-P., Planque, J., and Moutounet, M. (1999) Relative impact of major polysaccharides on the performance of an organic microﬁltration membrane. Am. J. Enol. Viticult., 50, 51. Gonçalves, F., Fernandes, C., and de Pinho, M.N. (2001) White wine clariﬁcation by micro/ultraﬁltration: effect of removed colloids in tartaric stability. Sep. Purif. Technol., 22–23, 423–429. Strathman, H. (1986) Electrodialysis, in Synthetic Membranes: Science, Engineering and Applications, NATO Asi Series C Mathematics and Chemical Science

18

19

20

21

22

(eds P.M. Bungay, H.K., Lonsdale and M.N. Pinho), Reidel Publishing Company. Escudier, J.-L., Moutounet, M., and Bernard, P. (1987) Filtration tangentielle. I. Inﬂuence de lultraﬁltratration sur la cinétique de cristallisation du bitatrate de potassium. Rev. Fr. Oenol., 108, 52. Jiao, B., Cassano, A., and Drioli, E. (2004) Recent advances on membrane processes for the concentration of fruit juices: A review. J. Food Eng., 63, 303–324. Vaillant, F., Jeanton, E., Dornier, M., OBrien, G.M., Reynes, M., and Decloux, M. (2001) Concentration of pation fruit juice on an industrial pilot scale using osmotic evaporation. J. Food Eng., 47, 195–202. Ferrarini, R., Versari, A., and Galassi, S. (2001) A preliminary comparison between nanoﬁltration and reverse osmosis membranes for grape juice treatment. J. Food Eng., 50, 113–116. Thang, V.H. and Novalin, S. (2006) Production of rectiﬁed grape musts by ion exclusion chromatography. J. Food Eng., 76, 128–213.

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6 New Applications for Membrane Technologies in Enology Martine Mietton Peuchot 6.1 Reduction of Alcohol Content

The development of techniques for reducing sugar content in musts and alcohol content in wines is the result of problems that certain vineyards are faced with, of overconcentration of sugars and, therefore, of alcohol in wines. The second reason for decreasing alcohol in the end product is that it corresponds to the wish of certain countries to take measures to restrain alcohol consumption. The problem is partly due, originally, to evolutions in viticultural and wine-making practices. It is now recognized that the quality of wine is a function of the phenolic compounds in the grape berries. Phenolic maturity is directly linked to a high concentration of sugars. This evolution in viticultural practices leads to too high a concentration of sugars that, in turn, can slow down or stop fermentation. A second cause of higher concentrations of sugar may be an increase in average temperatures due to evolutions of the climate. Knowledge of membrane technologies in the domains of reverse osmosis, nanoﬁltration, and ultraﬁltration has allowed the development of innovative alternatives in the partial reduction of alcohol content in wines [1–4]. Given that traditional methods of alcohol reduction are essentially based on phase changes (evaporation and distillation), there may be associated risks of deterioration of quality [5]. This has lead research teams and companies to introduce membrane technologies in alcoholreduction processes in order to propose solutions that are more selective and more respectful of the end product. 6.1.1 Reduction of Must Sugars to Obtain a Lower Alcohol Content in Wines

The process is based on a preventive action, reducing the sugar content of musts, thus allowing better control of subsequent alcoholic fermentation. This process is patented by the Bucher Vaslin company and is at present marketed under the name REDUXÒ [6, 7]. The heart of the system comprises two coupled membrane units that produce a concentrate, which represents the fraction to be eliminated. The

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Figure 6.1 Representation of the REDUXÒ process [7].

sugar-reducing process (Figure 6.1) consists of a ﬁrst stage, ultraﬁltration that produces a clear must of the same sugar concentration as the initial must. In a second stage, this must is concentrated by nanoﬁltration and the permeate that is produced, made up essentially of water and acids, is reincorporated into the must that is being treated. Nanoﬁltration gives greater ﬂow rates than reverse osmosis. The higher cut-off threshold of the membrane allows the transfer of acids and potassium into the permeate. This allows the partial reintroduction of acidity with the water recuperated from the must before fermentation, the advantage being that mature musts have, in general, low acidity levels and any elimination of this acidity would be detrimental. The acid balance of the treated must will be little affected by the process. The reduction in the sugar content of the musts, together with the loss of volume inherent in the treatment, leads to a modiﬁcation of the solids-to-liquid ratio during fermentation on skins, which in turn leads to an increase in phenolic compounds, tannins, anthocyanins, and potassium (Table 6.1). The highest sugar concentration in the concentrate obtainable by nanoﬁltration is in the region of 400 g l–1. However, differences in alcohol content during fermentation on the skins can modify the quantity and quality of the compounds that are extracted or produced. One notable phenomenon is that a lower alcohol content diminishes the burning sensation given by alcohol, with a corresponding reduction in sugariness and fatness. Nevertheless, this modiﬁcation of the organoleptic qualities of the wine does not mean that appraisals of wines made from a sugar-reduced must are any less favorable. The RAWÒ process, in which the concentration of ultraﬁltered must is achieved by evaporation in a vacuum instead of by nanoﬁltration, results in lower losses in volume, the sugar being eliminated in a more concentrated form. However, the water that is recuperated and then reincorporated into the musts is acid free, which can lead to reduced acidity in the must.

6.1 Reduction of Alcohol Content Table 6.1 Analytical results of wines from treated musts [7].

Merlot

Alcohol content (%.vol) Total acidity (TA) (gH2SO4 L1) Volatile acidity (VA) (gH2SO4 L1) pH Potassium (mg L1) IPT Tannins (g L1) Anthocyanins (mg L1) HCl index Dialysis index Gelatin index

Wine without treatment

Wine with ReduxÒ treatment

15.1 3.4 0.57 3.80 1420 77.3 4.7 654 31 30 40.8

14.2 3.5 0.58 3.85 1620 81.6 5.0 676 33 32 41.8

6.1.2 Reduction of Alcohol Content in Wine

In the 1980s a large number of studies were carried out on the partial or total removal of alcohol from wines. The reasoning was principally economic, given that wine growers the world over were suffering from chronic overproduction, and lesserquality wines were hard to sell. This is the background to the attempts that were made to develop new wine-based beverages or grape juices adapted to consumer tastes, and also to create new sales opportunities by producing new wine-based products such as light and alcohol-free wines. Unfortunately, the commercial success was, at the time, far from that expected. However, this subject is, at present, once more under study for the reasons stated above. The most commonly used alcohol-reduction technique is based on the selective separation of water and alcohol from the wine by reverse osmosis (the aromatic compounds are preserved) together with the separation of alcohol from the reverse osmosis (RO) permeate by distillation (D). Partial alcohol reduction is thus achieved without watering down, this being prohibited by a large number of wine-producing countries. This removal of alcohol is carried out after malo-lactic fermentation for red wines, and at the end of alcoholic fermentation for white wines. This process presents the advantage of being alcohol selective, but treatment capacities are limited: low ﬂow rates through reverse osmosis membranes, combined with low permeation rates, mean that it is necessary to work with large membrane areas and high pressures. This, of course, implies high investment and operating costs. Nanoﬁltration (NF), on the contrary, provides substantially higher alcohol ﬂow rates together with greater permeation rates. Working pressures are lower, leading to savings in investment and operating costs. In spite of a lower degree of selectivity in terms of the aromas contained in the wine, the organoleptic repercussions of

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Table 6.2 Analysis results of permeate and wine after removal of alcohol [7].

Alcohol content(% vol) pH Total acidity (g H2SO4 L1) Volatile acidity(g H2SO4 L1) Tartaric acid (g L1) Lactic acid (g L1) K þ (g L1) Anthocyanins (mg L1) IPT(Polyphenolic content, OD 280 nm)

Wine before NF

Permeate NF

Wine after NF

Wine after treatment

12.7 3.64 3.9 0.2 1.22 1.4 1.0 999 56.9

10.9 3.75 1.5 0.18 0.5 0.7 0.2 nm nm

13.1 3.65 4.1 0.2 1.4 1.6 1.2 1120 68.9

10.9 3.62 3.6 0.19 1.0 1.0 1.2 914 56

using nanoﬁltration for alcohol reduction are very close to results obtained by reverse osmosis [6]. Loss of aroma is compensated by the extraction of lower volumes in nanoﬁltration. Table 6.2 gives the results of analyses of wines and intermediate products carried out during partial alcohol removal in Cabernet Sauvignon wine. Alcohol removal does not signiﬁcantly modify the physical or chemical components of wine, only the acidity is slightly affected. Other membrane processes for alcohol reduction are also available. For example, the Australian process marketed by Memstar consists of alcohol reduction by a twostage process of nanoﬁltration followed by a membrane module known as a membrane contactor (Liqui-CelÒ ). The disadvantage of this process is that it uses large quantities of water to extract alcohol in the membrane contactor (MC). Table 6.3 compares different processes that are applicable to the reduction of alcohol in the end product. The sensorial evaluation has shown that, despite some aroma losses during the partial dealcoholization, the panel could not perceive some differences between initial wine of Merlot grape variety and the dealcoholized one. It can be concluded that this technology is feasible to achieve an alcohol reduction of 2% (v/v), without a perceptible depletion of the product quality [8].

Comparison of different processes for a reduction of 2% in alcohol content from a 14% wine [9]– RO: reverse osmosis; NF: nanofiltration; D: distillation; MC: membrane contactor.

Table 6.3

Volume of permeate/volume of wine (%) Volume of water (L) for the treatment/L wine Coproduct (efﬂuent)

RO-D

NF-D

RO-MC

NF-MC

25

18

50

30

0

0

0,45

0,3

Alcohol (92%)

Water with alcohol (4% RO, 7% NF)

6.2 Reduction of Malic Acid in Grape Musts or Volatile Acidity in Wines

6.2 Reduction of Malic Acid in Grape Musts or Volatile Acidity in Wines

The volatile acidity or malic-acid reduction could also be done by coupling two stages of reverse osmosis. Since the free acids are poorly retained by the membrane, the permeate after the ﬁrst stage ﬁltration (permeate 1) contains free acids, salts, esters and other small molecules. Once the permeate 1 is neutralized with pH of the targeted acid, it will be retained by the second stage membrane in a salty form. The other components passing through (permeate 2) are reinjected in the initial wine. Potassium hydroxide is used for neutralization. During the treatment, the decrease of the acid concentration in the must or wine is progressive. For example, the rejection rate of acetic acid in the ﬁrst reverse osmosis system varies between 40 and 50%. After the neutralization, the rejection rate of potassium acetate in the second reverse osmosis is higher than 90%. 6.2.1 Reduction of Malic Acid in Musts

Different methods can be used to remove acids from grape must or wine: cold stabilizing (partial precipitation of potassium tartrate), by the addition of chemicals such as calcium carbonate, or by malo-lactic fermentation [10, 12]. During malo-lactic fermentation, the transformation of 1 g l1 of malic acid into lactic acid results in a reduction in acidity corresponding to 0.6 g l1 of tartaric equivalent. Moreover, the microbiological stability of the wine is enhanced. However, malo-lactic fermentation is not always easy to control and it may provoke signiﬁcant changes in the aromatic proﬁle of wines, such as an increase in lacteous and buttery characteristics or a decrease in fruitiness [123]. The membrane process for the removal of malic acid is performed in two nanoﬁltration stages. The racked must is nanoﬁltered. The permeate contains water, malic acid, tartaric acid, and traces of small constituents contained in the must (Figure 6.2). The nanoﬁltration permeate is neutralized to a pH of approximately 7 by using potassium hydroxide. C4 H6 O52 þ 2 KOH ! ½C4 H4 O5 2 þ 2 K þ þ 2 H2 O

In the second stage, the neutralized permeate is nanoﬁltered through the same membrane. The potassium malate is thus retained by the membrane. The permeate is reincorporated into the must. For a continuous process using two membrane units, the permeate ﬂow rates of the two membranes should be identical in order to allow correct control of the pH during neutralization. This can be achieved by adjusting operating pressure. 6.2.2 Reduction of Volatile Acidity

The winemaking process can sometimes produce volatile compounds that impair the quality of the wine. Among these undesirable compounds, the best known is acetic

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Figure 6.2 Influence of pH on the composition of the second permeate [13].

acid, present in free form or as a salt or ester. Processes have been developed to reduce volatile acidity in wines. They combine a nanoﬁltration or reverse osmosis stage with ion-exchange resins: following reverse osmosis of a fraction of the wine, the resulting permeate containing acetic acid is subjected to treatment on weak anionic resins and then reincorporated into the initial wine (Vinovation process). The process depends on the ability of a semipermeable membrane to separate from wine a permeate stream containing acetic acid and ethyl acetate, but substantially no ﬂavor or color. Wine from the tank is recirculated via tangential ﬂow against a reverse-osmosis membrane, and a small portion passes through. The retentate contains all the ﬂavor and color, and is returned to the tank. The permeate is a colorless, ﬂavorless liquid containing only water, alcohol, acetic acid and ethyl acetate, and is totally devoid of vinous character. The permeate is passed through a weak-base anion-exchange resin. Ethyl acetate is hydrolysed by the basic conditions within the column. The resin retains acetic acid, while permitting alcohol and water to pass through. The puriﬁed permeate is then recombined with the retentate and returned to the tank. The process continues until the desired degree of volatile acidity reduction is achieved. The resulting wine is essentially unchanged in volume and ﬂavor. The main disadvantage of this method arises from the use of resins, which require regular regeneration. The second process that is proposed is a combination of two membrane processes. It is based on the fact that reverse osmosis and nanoﬁltration membranes have different retention properties according to pH. Thus, in the case of a weak acid, the membrane will allow compounds with a low pH to pass through and will have a high rate of rejection above its pK. Acetic acid in wine having a pH lower than the membrane pK (4.75), it will not be retained. In a saliﬁed form, at a pH higher than that of the pK, it will be retained. The ﬁrst stage of the process consists of reverse osmosis of the wine, giving a permeate that is relatively rich in acetic acid. This permeate, neutralized by potassium hydroxide, is subjected to osmosis in the second stage of the process. The potassium acetate is retained by the membrane and the acid-reduced

6.3 Acidification of Musts and Wines

permeate is then reincorporated into the partially concentrated wine, which thus recovers a normal level of volatile acidity. The ﬁrst reverse osmosis allows 50% of the acetic acid to pass through, whereas more than 99% of malic and tartaric acids are retained (compounds that are more voluminous and pK2 higher than that of wine). The retention rates for acetic acid and potassium from the permeate, for different volatile-acidity neutralization levels, are determined by the use of reverse osmosis and nanoﬁltration membranes. Both types of membrane retain over 98% of potassium acetate at a pH of 10.

6.3 Acidification of Musts and Wines

Over the last ﬁfteen years or so it has been observed that wines have been appearing that contain higher ethanol levels and pH values that are higher and higher. This phenomenon has been attributed to global warming. Present pH values range from 2.8 to 4.2. This gradual rise in the pH values of wine generates problems in terms of the control of the evolution from both the microbiological point of view and that of color stability. This situation has led to greater and greater use of tartaric acid to acidify wines. Acidiﬁcation ought to contribute to the balance of the gustative sensations provided by wines, promote correct biological development as well as adequate conservation of the wine, and compensate for a lack of natural acidity caused by local climate conditions or by wine-making practices that result in a lowering of natural acidity. Taking into account the saliﬁcation balance of the organic acids in wine, the acidifying effect will result from a reduction in the proportion of saliﬁed forms and, thus, in mineral cation content. The aim of acidiﬁcation is to modify pH values and not titratable acidity, which means increasing the proportion of free acids to the detriment of saliﬁed forms [14]. Conventional electrodialysis techniques have been successfully used in enology for the tartaric stabilization of wines [15]. The bipolar membrane is a thin polymer wall that is rendered operational by ion-exchange layers. The bipolar membrane has an anion-exchange face and a cation-exchange face, and functions in the same way as separate anionic and cationic membranes. The role of the bipolar membrane is to maintain the acid/base ionic balance of the process, this being achieved by the electrolysis of water molecules in the bipolar membrane under the driving force of an electric ﬁeld during the treatment. The bipolar membrane must be correctly oriented: the cation-exchange side facing the cathode is permeable only to cations. In this way, a stack of bipolar membranes with cationexchange membranes will only allow the passage of cations while retaining anions as well as uncharged particles [16]. This operation causes acidiﬁcation by lowering the pH. The stacking of bipolar membranes in association with cation-exchange membranes means that there are two parallel hydraulic circuits: the compartment called the diluate, which contains the wine, is acidiﬁed, whereas the concentrate compartment, containing an ionic solution, becomes more alkaline. The electric current that is applied between the two electrodes splits water molecules into OH and H þ inside the bipolar membrane, which is in contact with the wine. The OH

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ions migrate towards the positive pole (anode) into the brine, whereas the H þ ions migrate towards the negative pole (cathode) and make up for the potassium ions that are extracted from the wine. When the electric current is applied, the potassium ions (K þ ) contained in the wine are attracted towards the cathode, they pass through the cationic membrane and are stopped by the bipolar membrane. The H þ ions passing through the bipolar membrane then replace the K þ ions in the wine in order to conserve the ion equilibrium. Acidiﬁcation by bipolar electrodialysis can correct wine pH with a precision of 0.05 units. The target value of the treatment is determined following tasting with the producer. The maximum treatment value is 0.3 units of pH. Following acidiﬁcation, malo-lactic fermentation takes place in the different wines, leading to slight variations in the pH values obtained. From the organoleptic point of view, products treated by electrodialysis are perceived as being fresh and not so heavy in the mouth. The phenomenon of harshness in the mouth, which is the main disadvantage of the addition of tartaric acid, is not remarked upon, and color is also more intense [15]. Lowering the pH of wines essentially means a reduction in the concentration of potassium. The anion content remains unchanged, as is shown by the analyses of the organic acids and the determination of volatile acidity levels. For a lowering of pH values there is a concomitant increase in titratable acidity. Other analytical criteria (such as must sugars, alcohol content, residual sugars in wine, etc.) are not affected by the treatment, which only concerns positively charged elements. The product is treated in a continuous process controlled via on-line readings of pH values. The process can be fully automated and requires only one treatment cycle, with no need for recycling [15]. Real-time supervision allows decisions to be taken at the right moment as a function of reasoned technological or commercial objectives. The treatment is carried out at normal temperature and atmospheric pressure with no mixing or stirring, through a series of membranes until the desired pH is obtained. Treatment time is short; the membrane modules are relatively small and efﬁcient: this new technology is, therefore suitable for installation in mobile units so that it can be made widely available on a subcontract basis. The low-pollutant waste products can be easily used, for example for cleaning, and need not penalize the process.

6.4 Other Potential Applications

The reduction of bad tastes seems to be one of the ﬁelds being at present explored. Chilean researchers [17] have been working on the reduction of 4-ethylphenol and 4ethylguaiacol in red wines by nanoﬁltration and adsorption. The permeate from the membrane is put into contact with a hydrophobic adsorbant resin, XAD-16HP. This is circulated in the wine until the desired concentrations are obtained. The process also allows the elimination of herbaceous aromas. The results of the study show that the resin is not sufﬁciently selective. The isolation of bad tastes in nanoﬁltration

References

permeate is possible, but it is necessary to ﬁnd a way to eliminate theses bad tastes by a speciﬁc treatment (adsorption, ﬁning, etc.).

References 1 Wucherpfennig, K. (1978) Possibilit es

2

3

4

5

6

7

8

9

dapplication de procedes utilisant des membranes pour la stabilisation des vins (ultraﬁltration, osmose inverse, electrodialyse). Ann. Technol. Agric., 27 (1), 319–331. Berger, J.L. (1991) Auto-enrichissement du moût par osmose inverse. Bull. OIV., 721–722, 189–210. Cuenat, P., Kobel, D., and Zuffrey, E. (1989) Losmose inverse et lœnologie. Bull. OIV., 701–702, 519–537. Mietton-Peuchot, M., Milisic, M., and Noilet, P. (2002) Grape must concentration by using reverse osmosis. Comparison with chaptalization. Desalination, 148, 125–129. Escudier, J.L. (1989) Evaluation de quelques techniques adaptees a la desalcoolisation des vins. Rev. Fr. œno., 119, 56–61. Cottereau, P., Solanet, D., Vuchot, P., Ferment, E., and Noilet, P. (2006) Reduction de la teneur en sucre des moûts ou de la teneur en alcool des vins. Actes du congres de lOIV 2006, p. 9. Noilet, P. (2004) Contribution aux differentes applications de losmose inverse dans le traitement des moûts et des vins. Universite Bordeaux 2, Diplôme Universitaire dExperimentation et de Recherche en Œnologie et Ampelologie: 135p. Diban, N., Athes, V., Bes, M., and Souchon, I. (2008) Ethanol and aroma compounds transfer study for partial dealcoholization of wine using membrane contactor. J. Membrane Sci., 311, 136–146. Aguera, E., Athes-Dutour, V., Bes, M., Caille, S., Cottereau, P., Escudier, J.L., Mikolajczak, M., Roy, A., Sablayrolles,

10

11

12

13

14 15

16

17

J.M., Samson, A., Souchon, I., and Vidal, J.P. (2009) Reduction de la teneur en alcool des vins - Etude comparative de differentes technologies. Symposium OIV, 28 June–3 July, Zagreb, p. 8. Gerland, C. and Poinsaut, P. (2002) Pratiques oenologiques et acidite des vins. Revue des OEnologues, 105, 19–22. Ducruet, J. (1997) Incidence du mode de logement lors de la fermentation malolactique sur la qualite des vins rouges. J Sci Techn Tonnellerie, 3, 61–66. Dols-Lafargue, M., Gindreau, E., Le Marrec, C., Chambat, G., Heyraud, A., and Lonvaud-Funel, A. (2007) Changes in red wine soluble polysaccharide composition induced by malolactic fermentation. J. Agric. Food Chem., 55, 9592–9599. Ducruet, J., Fast-Merlier, K.L., and Noilet, P. (2010) New application for nanoﬁltration: reduction of malic acid in grape must. Am. J. Enol. Viticult., 61, 2. Usseglio-Tomasset, L. (1989) Chimie Oenologique, Lavoisier, Paris. Moutounet, M., Saint-Pierre, B., and Batlle, J.L. (2005) Acidiﬁcation des moûts et des vins par un Procede Electromembranaire a Membranes Bipolaires. Meeting of expert group of OIV. Bazinet, L., Lamarche, F., and Ippersiel, D. (1998) Bi-polar-membrane electrodialysis: Applications of electrodialysis in the food industry. Trends Food Sci. Tech., 9 (3), 107–113. Ugarte, P., Agosin, E., Bordeu, E., and Villalobos, J.I. (2005) Reduction of 4-Ethylphenol and 4-Ethylguaiacol concentration in red wines using reverse osmosis and adsorption. Am. J. Enol. Vitic., 56 (1), 30–36.

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7 Membrane Emulsification for Food Applications Henelyta S. Ribeiro, Jo J. M. Janssen, Isao Kobayashi, and Mitsutoshi Nakajima 7.1 Introduction

Membrane emulsiﬁcation has attracted increasing attention to pharmaceutical, chemical, food and cosmetic industries in the last decade. As an innovative process it may provide reduction in energy, chemicals consumption and waste production [1]. Operational ﬂexibility and reduction in the ratio of equipment size to production capacity, easy scale-up, and reproducibility may lead membrane emulsiﬁcation technologies to considerable efﬁciency. It may be suitable for industrial-scale production as a novel process intensiﬁcation, eliminating major large-scale equipments that require high maintenance costs. Nowadays, the food industry is putting considerable effort into the manufacturing of products with high quality, nutritional value and a natural taste. Appropriate processing methods are at the core of this development, because processing determines the product microstructure to a signiﬁcant extent. Moreover, delicate ingredients and structural elements can be adversely affected in their functionality and nutritional value if the processing is too harsh. In the past decades, membrane emulsiﬁcation (ME) has been identiﬁed as a promising method for making single and multiple emulsions, solid lipid colloids, gel and core shell particles under relatively mild conditions [2–9]. Some recent literature has also reported the production of nano- and microbubbles using cross-ﬂow membrane emulsiﬁcation [10–14]. Several industries have been investing in the development of these technologies, which may lead to a new process route and equipment [4, 13, 15–23]. ME and microchannel emulsiﬁcation (MCE) are low energy input process (103 to 106 J m3), and have the potential to produce very narrow droplet-size distribution compared to other emulsiﬁcation techniques, with special application to parenteral emulsions as droplet-size distribution can be easily controlled by process parameters. Experimental studies and modeling analysis have shown a great advance on the understanding of droplet formation and its uniform droplet-size distribution using membrane (ME) and microchannel (MC) emulsiﬁcation processes [19, 24–33], which will be highlighted in Section 7.2.

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Irregular microstructure and surface properties of current membranes still limit the full exploitation of the beneﬁts of ME, where its major disadvantage is the low ﬂow rate compared to conventional mechanical emulsiﬁcation processes. Also membrane fouling by particulates or adsorbing species can be an important problem. Particulate fouling will block pores, while adsorbing species (which may even be emulsiﬁers from the product formulation) can change the wetting properties of the membrane. Various membrane emulsiﬁcation technologies and membrane materials have been developed; current operating methods include cross-ﬂow (XME), dead-end (PME), rotating (RME), and vibrating (VME) membrane emulsiﬁcation, as well as microchannel emulsiﬁcation (MCE). This chapter will describe each of these methods and their applications. 7.1.1 Cross-Flow Membrane Emulsification (XME)

XME is considered the conventional membrane emulsiﬁcation process to prepare uniform droplets of oil-in-water (O/W) and water-in-oil (W/O) emulsions, and also multiple emulsions. Important process parameters to be considered in XME are: transmembrane pressure; type of membrane, permeability, and thickness; membrane pore size, porosity and wettability; type of emulsiﬁer and its concentration; cross-ﬂow and continuous phase velocity; viscosity of dispersed and continuous phases. In this process a relative low pressure forces the disperse phase through the membrane pores and the droplets detach from the pore outlets into the continuous phase containing an emulsiﬁer for immediate droplet stabilization. Interfacial tension, inertial, buoyancy, dynamic lift, drag and static pressure difference forces act on a droplet during this membrane emulsiﬁcation process. It is necessary to have a balance between all these forces for the success of the droplet formation [3, 9, 34–36]. Under model conditions the process can be monitored by a high-speed camera and optical microscopy. For the production of O/W and W/O emulsions, either hydrophilic or hydrophobic membranes, respectively, are required. Examples of hydrophilic membranes include ceramic (a-Al 2O3, zirconium oxide) and metallic, microporous glass membrane made of calcium aluminoborosilicate glass synthesized from a volcanic ash called Shirasu (CaO-Al2O3-B2O3-SiO2, Shirasu porous glass, SPG), polypropylene, polycarbonate, polyvinylidene ﬂuoride (PVDF), and poly(tetraﬂuoroethylene) (PTFE) [37]. They can be made hydrophobic by chemical surface modiﬁcation, especially organic silane coupling agents (e.g., octadecyltrichlorosilane), which are nonfood grade. A range of membrane with pore size ranging from 0.05 to 30 mm is commercially available. The pore diameter of the membrane is the crucial parameter to determine the ﬁnal droplet size of the emulsion and the distance between two adjacent pores should be far enough to prevent coalescence of forming droplets. Wettability of the membrane is very important for the process performance, allowing narrow droplet-size distributions [2]. Membranes should always be wetted by the continuous phase before starting the emulsiﬁcation

7.1 Introduction

process. To maintain membrane performance, it is important to avoid wetting by the disperse phase. More recently, Kukizaki [21] has used SPG membranes to study the droplet formation behavior in the absence of shear ﬂow at the membrane surface. A faster decrease in the interfacial tension and slightly higher transmembrane pressure than the capillary pressure allowed spontaneous formation of smaller droplets with narrower size distribution. 7.1.2 Dead-End Membrane Emulsification (PME)

Suzuki et al. [38, 39] have reported the ﬁrst research work on the application of PME to produce O/W and W/O emulsions. In this membrane emulsiﬁcation technology a preliminary coarse emulsion is forced through the porous membrane and mother droplets are broken up into daughter droplets resulting in smaller droplet sizes and narrower droplet-size distributions than the pre-existing emulsion. Higher transmembrane ﬂux and easier operational conditions makes PME more advantageous than XME for large scales, however wider droplet-size distribution can be observed. For the production of more monodisperse emulsions, a number of passes of the emulsion through the membrane is required, which make the process more expensive than XME. Different types of membrane materials have been used in PME, SPG membranes are the most conventional ones. Some authors have also used PTFE [40, 41], polyamid 6,6 [42, 43]. Increase of the disperse phase fraction results in lower transmembrane pressure for any type of membrane used. As in XME, different surface properties of the membranes provide the production of either O/W or W/O emulsions (hydrophylic and hydrophobic membrane, respectively). Vladisavljevic et al. [44] investigated the inﬂuence of anionic, nonionic and zwitterionic emulsiﬁers on the mean droplet size, transmembrane ﬂux, and membrane fouling in repeated PME using SPG membrane. Control of pH may allow better performance during processing, mainly when protein is used as the emulsiﬁer as agglomeration occurs at pHs closed to its isoelectric point, resulting in strong membrane fouling and low transmembrane ﬂux. 7.1.3 Rotating Membrane Emulsification (RME)

Rotating disk/cylindrical membranes have been applied to dynamic membrane ﬁltration in large scales. The success of this ﬁltration method was the motivation for the development of the novel rotating-membrane emulsiﬁcation process. More recently some research groups have shown interest in its application due to increasing in ﬂow rate of the dispersed phase through the membrane, which single and multiple monodisperse emulsions could be successfully produced [45–49]. In this technique, shear stress is developed by rotating a cylindrical membrane and disperse phase is radially forced through its pores into the continuous phase

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containing the emulsiﬁer, allowing droplet stabilization. Vladisavljevic and Williams [48] studied the production of O/W emulsions using RME at different angular speeds (50–1500 rpm). A stainless steel membrane with pore diameter of 100 mm was used, and showed the best performance at 350 rpm where droplet sizes up to 107 mm (coefﬁcient of variation, CV ¼ 4.9%) were produced. The literature has shown a consistent behavior of the RME process where the angular speed significantly inﬂuences the size of the droplets formed. Droplet size tends to decrease at higher angular speeds [48]. RME is a potential process to be applied in industrial scale, mainly for the production of larger droplet size [47]. Further development is required since it is a relative new membrane emulsiﬁcation process 7.1.4 Vibrating Membrane Emulsification (VME)

In the previous section, membrane rotation was discussed as an alternative to crossﬂow for creating a shear force on the droplets that form at the membrane pores. Membrane vibration is another option, and may be more appropriate for ﬂat membranes like perforated metal plates and microengineered silicon wafers. Zhu and Barrow [50] have studied the effect of lateral piezoactuated vibration of thin microengineered silicon nitride membranes in a laboratory-scale rig. The formation of droplets at individual pores (with large interpore distance) was observed with videomicroscopy. In order to enable these observations the timescale of droplet formation could not be chosen too short, so crossﬂow and dispersed phase ﬂow had to be kept rather small. This resulted in rather large droplets for the stationary membrane case, that is, of the order of 100 mm. Membrane vibration was observed to reduce the droplet size, but only at rather low frequency and not to a very large extent. Kelder et al. [19] have studied XME with vibrating membranes theoretically, using a simple analytical force-balance model as well as 3D computational ﬂuid dynamics (CFD) simulations. These authors showed that the effect of membrane vibration on droplet formation is quite complex. First, the drag force due to the vibration should be at least comparable to the drag exerted by the crossﬂow in order to have a signiﬁcant effect. Moreover, the frequency of the vibration should be linked to the droplet-formation frequency in order to assure that each droplet is affected by the vibration in the same way. Outside this resonance the droplet behavior can become quite irregular, leading to a wide size distribution. Kelder et al. [19] also considered the power requirements for the vibration. The average power was found to be proportional to the square of the vibration amplitude and to the third power of the frequency. The power input m3 of emulsion was estimated as well, and was found to be of the order 105 W m3 for typical conditions. This is signiﬁcant compared to the overall energy input and diminishes the energyefﬁciency advantage that is usually attributed to XME. Overall, one can conclude that membrane vibration is not an obvious option for large-scale applications, and that even for small-scale specialty applications considerable technical challenges still have to be overcome.

7.1 Introduction

7.1.5 Microchannel Emulsification

Several types of ME have been introduced in the previous sections. The smallest droplet-size distribution for ME is approximately 10% in CV, primarily due to poresize distribution and/or sensitivity to operating conditions. Monodisperse emulsions consisting of highly uniform droplets have recently received great attention in various ﬁelds including foods, pharmaceuticals, cosmetics, and chemicals. Nakajima and colleagues proposed microchannel emulsiﬁcation (MCE) for producing highly uniform droplets with a small coefﬁcient of variation of below 5% in the 1990s [51]. The droplet generation unit (DGU) used in MCE is a microchannel (MC) array consisting of parallel MCs with a terrace and a deep well (Figure 7.1a). Droplets are directly generated in the well via an MC array, even in the absence of a cross-ﬂowing continuous phase. This droplet generation based on spontaneous transformation is a very mild process and has very high energy efﬁciency (e.g., 65% in Sugiura et al. [25]). Dead-end MCE chips were used in the initial stage of MCE researches, with Nakajima and colleagues applying MC array chips developed for analyzing blood rheology in blood capillaries [52] to emulsiﬁcation. Interestingly, the terrace, which plays an important role in droplet generation by MCE, was originally designed for observing the behavior of blood components. Kawakatsu et al. [53] designed cross-ﬂow MCE chips for long-term operation and emulsion collection. However, MCE chips consisting of grooved MC arrays (Figure 7.1a) have a very low productivity of vegetable oil droplets (<1 L m2 h1) when MCs with a size of 10 mm are used. A straight-through MC array consisting of highly integrated microﬂuidic through-holes (Figure 7.1b) remarkably improved the droplet productivity of MCE [54]. Straight-through MC arrays with an MC size of 10 mm produced uniform vegetable oil droplets at a maximum dispersed-phase ﬂux of 60 L m2 h1 [55]. Currently, MCE is capable of producing monodisperse emulsions with a droplet

Figure 7.1 Schematic drawings of droplet generation via part of a grooved MC array (a) [51] and part of a straight-through MC array (b) [83].

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size of 1 to 100 mm [56, 57]. Monodisperse emulsions produced by MCE also have been used as templates for obtaining monodisperse microdispersions such as microparticles and microcapsules. Later sections will discuss the process fundamentals of MCE and the production of emulsions and microdispersions for food applications using MCE.

7.2 Understanding of the Process at the Pore Level

The literature on droplet formation in ME and MCE is extensive, and a detailed discussion is beyond our current scope. Rather, we intend to provide a short overview of the current understanding and recent developments, referring to key papers for further details and additional references. We will subsequently discuss XME, PME and MCE. 7.2.1 XME

The growth and detachment of a droplet at a (often circular) pore in a cross-ﬂow has been studied in detail over the last decade using high-speed videomicroscopy, computational ﬂuid dynamics (CFD), surface evolver and lattice-Boltzmann (LB) simulations [19, 27, 33]. These studies show that the growing droplet is initially displaced from the pore in the direction of the cross-ﬂow, while remaining attached to it by a thin neck. When the growing droplet has become too large for the neck to resist the drag force on the droplet, the neck ruptures and leaves part of its volume attached to the pore as the starting point for the formation of the next droplet. Kelder et al. [19] noted in their CFD simulations that the rupture of the neck occurs close to the droplet rather than close to the pore, and hypothesized that the part of the neck close to the pore is stabilized by the centrifugal force that is due to the bending of the droplet phase ﬂow direction over almost 90 . Further down the neck this stabilizing effect has disappeared and instability can occur more easily. These and other studies have provided some quite useful insight, but it is worth noting that for experimental and computational reasons the details of the neck formation and behavior have been studied under conditions where rather large droplets are produced. It is not clear at present if all the details translate completely to the droplet size range (well) below 20 mm, which is typical for many food emulsions. For process optimization and scale-up, the simple mechanical models that were introduced ﬁrst by Schr€oder et al. [58] and Peng and Williams [24] are still preferred. These authors (and many after them) have shown that a number of forces act on the forming droplet, but that several of these are usually orders of magnitude too small to be relevant. Peng and Williams [24] have retained only the two largest contributions, that is, the drag exerted on the forming droplet by the cross-ﬂow and the interfacial tension force that keeps the droplet attached to the pore. Their ﬁrst model was a torque balance, as shown at the lefthand side of Figure 7.2. In this model the droplet is

7.2 Understanding of the Process at the Pore Level

Figure 7.2 Torque balance and force-balance models, after Peng and Williams [24].

assumed to grow as a hemispherical cap on top of the pore. Peng and Williams [24] have shown that the ﬁnal droplet radius RD is then given by: 3 RD s ¼ ð7:1Þ RP 5:1 tW RP Here, tW is the wall shear stress, which is given by: tW ¼

1 r f W2 2

W is the relevant mean velocity, r is the density and f is the friction factor, which depends on the channel Reynolds number. It is noted that this torque model does not invoke the notion of a neck. Peng and Williams [24] also brieﬂy introduced a simple force balance, which does invoke the notion of a neck between the droplet and the pore. Near the critical conditions for detachment the neck is supposed to have bent about 90 towards the direction of the cross-ﬂow, while more or less retaining its radius, as sketched at the right-hand side of Figure 7.2. The interfacial tension force is then comparable in magnitude to the torque case, but directed opposite to the drag force (hence a force balance rather than a torque balance). The ﬁnal equation of this model is very similar to that of the torque balance, that is, the exponent 3 is merely replaced by an exponent 2. Kelder et al. [19] have shown that this simple force-balance model agrees fairly well with CFD simulations in which neck formation is observed. It is noted that De Luca et al. [59, 60] have recently developed a different force-balance model, focusing on the forces that act on the contact line on the membrane surface and without invoking the formation of a neck. The above models refer to cases where the interfacial tension force and the drag force are dominant compared to the other forces. However, when the cross-ﬂow velocity is reduced to (almost) zero the hydrodynamic force exerted by the liquid that ﬂows into the droplet becomes the dominant force that leads to detachment. This regime, which is sometimes referred to as spontaneous detachment, has been studied in detail at the University of Soﬁa [61, 62]. In the ﬁrst paper these authors presented a detailed analysis of the hydrodynamic force exerted on the forming droplet by the liquid ﬂowing into it. In the second paper they made an analogy between the detachment of a droplet from a pore in zero cross-ﬂow and the gravity-induced detachment of a pending drop. In particular, they took the

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well-known theory for gravity-induced detachment of a pending drop and replaced gravity by the previously established hydrodynamic force exerted by the liquid. In this way they were able to provide a consistent quantitative analysis of the droplet formation in (almost) zero cross-ﬂow. Discussing details goes beyond the scope of this chapter, but we do note that this work explains for the ﬁrst time why many authors have observed a correlation RC/RP 3 over a wide range of pore sizes under conditions of transmembrane pressures just above the critical value and relatively small cross-ﬂow. It is frequently suggested in the literature that a correlation of the type RC/RP ¼ constant is inherent to XME, but the wide range of reported constants is then difﬁcult to understand. Moreover, one can already see in Equation 7.1 (in which a factor 1/RP also appears at the right-hand side) that this linear proportionality does not hold in the presence of a non-negligible crossﬂow. De Luca et al. [59, 60, 63] tried to modify and extend the mechanical models such that they do predict a linear correlation, but convincing results were not obtained. These authors who have reported a correlation RC/RP ¼ constant with a value for the constant that lies well above 3 have probably merely made a linear ﬁt to a limited range of pore sizes. Another extreme case is that in which the droplet phase ﬂows out of the pore so fast that it initially becomes a jet, which breaks up into fragments at some distance from the pore. This is a well-known phenomenon for macroscopic nozzles and oriﬁces, and Christov et al. [62] have shown it experimentally for liquid ﬂowing out of a thin (180 mm diameter) capillary. Lambrich and Schubert [35] and others have mentioned the jetting regime in ME as well, and Lambrich et al. [64] have presented experimental results on XME in the jetting regime for microengineered membranes. As pointed out by Lambrich and Schubert [35] the main advantage of operating in the jetting regime is the large droplet-phase ﬂux as compared to conventional XME conditions, but this can only be obtained at reasonable transmembrane pressure if (nearly) all pores can be made to produce jets. The hydrodynamic resistance per membrane channel then has to be rather small, which is fairly easy to realize for the thin microengineered membranes used by Lambrich et al. [64]. For a much thicker ceramic or SPG membrane this is probably not possible. Nevertheless, jetting can play a role for these membranes, as argued by Christov et al. [62], but this is seen in the droplet-size distribution rather than the overall ﬂux. The proposed mechanism is that the large interconnectivity of channels within a ceramic or SPG membrane allows for the possibility that many internal channels jointly feed a given pore at the membrane surface, which then gets a high exit velocity. Up to now we have discussed droplet formation at a single pore, neglecting interactions with neighboring pores. Whether or not these are important depends on the interpore distance (or membrane porosity). If pores are closer to each other than the critical diameter for a single pore, growing droplets can coalesce when not properly stabilized by emulsiﬁers or can push each other off the pore when well stabilized. The latter behavior has been observed among others by Zhu and Barrow [49] Egidi et al. [65] and Kosvintsev et al. [66]. Kosvintsev et al. [66] have also developed a mechanical model for the push-off effect for the case of zero cross-ﬂow. Typically this push-off effect leads to smaller droplets, since

7.2 Understanding of the Process at the Pore Level

growing droplets then cause each others detachment prematurely compared to the single-pore case. Another aspect that has received considerable attention in recent years is that of interfacial rheology, particularly the role of a dynamic interfacial tension. The formation of a droplet implies the creation of new interface, to which emulsiﬁers will adsorb. If the expansion rate of the droplet interface is large, the emulsiﬁer transport will not be able to keep up and the dynamic interfacial tension that determines droplet growth and detachment will be close to the value for a clean interface. By contrast, if the interfacial expansion rate is small the dynamic interfacial tension will be close to the equilibrium value. Schr€ oder et al. [58] were among the ﬁrst to consider these aspects, showing that rapidly adsorbing small-molecule emulsiﬁers produce smaller droplets than more slowly adsorbing macromolecules. Rayner et al. [67] and Van der Graaf et al. [27, 68] have considered this further. De Luca et al. [63] have discussed the incorporation of the dynamic interfacial tension into the torque- and force-balance models of Peng and Williams [24], which then require numerical solution. 7.2.2 PME

Premix membrane emulsiﬁcation (also known as dead-end ME) was introduced by Suzuki et al. [37, 38] as mentioned previously, and has since been studied by several authors. In the context of understanding the process at the pore level one can state that the modeling of PME has not yet been developed in much detail. One reason is the fact that experimental observations at the pore scale are lacking, since the droplet break-up behavior within the membrane cannot be observed with methods like videomicroscopy. Van der Zwan et al. [29] have recently tried to bridge this gap by monitoring droplet behavior in thin microengineered model structures between glass plates using video-microscopy, and they observed quite complex behavior. First, they noted that an accumulation of droplets within the model membrane occurred in all cases, probably because the transport of droplets through the membrane is hindered more by the internal structure (e.g., bending and diameter variations of the channels) than the ﬂow of the continuous phase. The behavior of individual droplets is thus greatly affected by the presence and behavior of neighboring droplets. Moreover, droplets can temporarily (or sometimes permanently) block certain channels. This causes rather erratic changes in ﬂow throughout the structure. Van der Zwan et al. [29] were able to distinguish three modes of droplet break-up, that is, break-up due to localized shear forces, break-up due to interfacial tension effects and break-up due to steric hindrance between droplets. Moreover, they observed that break-up also occurs outside the membrane, within the layer of accumulated droplets at the upstream side. Modeling of these phenomena to a similar level of detail as for XME is clearly very complex, and has not been attempted up to now. In subsequent work Van der Zwan et al. [69] used a bed of small beads as a model PME membrane, and found that a correlation between droplet size and energy input per unit volume can be established.

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7.2.3 MCE

Droplet generation by MCE has been investigated in detail using high-speed videomicroscopy, CFD, and LB simulations [25, 26, 30, 31]. Screenshots portraying droplet generation via an MC and a terrace are presented in Figure 7.3. The dispersed phase that passes through the MC gradually expands on the terrace (Figures 7.3a and b), and then the dispersed phase that passes through the terrace outlet starts to expand into a well. In the initial stage of this detachment process, the Laplace pressure of the dispersed phase on the terrace (dPLap,terrace) is lower than that in the well (dPLap,well). dPLap,well gradually decreases with the increasing size of the expanding dispersed phase in the well, whereas dPLap,terrace is almost constant during this stage. Afterwards, dPLap,terrace becomes signiﬁcantly higher than dPLap,well, causing rapid ﬂow of the dispersed phase into the well. In this case, the dispersed phase on the terrace shrinks rapidly until a neck is formed on the terrace (Figures 7.3c–e). This behavior is driven by interfacial tension [25]. When the

Figure 7.3 Screenshots of droplet generation for MCE. Refined soybean oil was used as the dispersed phase, and Milli-Q water containing 1 wt% sodium dodecyl surface (SDS), as the continuous phase.

7.2 Understanding of the Process at the Pore Level

dispersed-phase ﬂux at the neck exceeds that in front of the neck [31], the neck instantaneously pinches off and a droplet is generated (Figures 7.3e and f). During this process, the dispersed-phase pressure at the neck becomes remarkably higher than that on the terrace and in the well [26]. Uniform droplets are periodically generated by spontaneous transformation of the dispersed phase that passes through the MC in the absence of a cross-ﬂowing continuous phase. Given the dispersedphase ﬂow in a 10-mm MC, the interfacial tension is by several orders of magnitude greater than the gravitational force, inertial force, and viscous force [70]. The effect of interfacial tension becomes more dominant as the MC becomes smaller, which is advantageous for producing many food emulsions. Although key points of the MCE process have been clariﬁed in the literature, further work should be conducted to obtain a complete understanding of the droplet-generation process. In MCE, the droplet size is determined primarily by the geometry of the MC array and can be tuned by changing the viscosity ratio of the two phases. Sugiura et al. [71, 72] studied the effect of the MC and terrace dimensions on droplet size. The parameters most affecting the resultant droplet diameter were the MC (and terrace) depth and the terrace length [71]. Analytical models for predicting the droplet size for MCE have been proposed by van Dijke et al. [31] and Sugiura et al. [71]. These prediction models consider the effects of the MC and terrace structures, the dispersed-phase pressure, and the interfacial tension and contact angle, but do not include the viscosity effect. The MC width and length, which hardly affect the droplet size, are the parameters affecting the droplet productivity per MC [72]. Using long, square MCs leads to generation of uniform droplets of a speciﬁc size at high productivity due to the great pressure drop of the dispersed phase in the MC. The droplet size is not sensitive to the ﬂow velocity of the dispersed phase inside the MC or to the applied pressure of the dispersed phase below a critical value, unlike XME and RME. This robust feature is advantageous for the practical production of monodisperse emulsions. Sugiura et al. [73] investigated the ﬂow state of the dispersed phase during MCE and reported that the character of droplet generation from MCs is determined by a dimensionless number called the capillary number (Ca), deﬁned as the ratio of the viscous force to the interfacial tension. The reported critical Ca was approximately 0.02, indicating that the interfacial tension basically dominates the ﬂow state of the dispersed phase during droplet generation. Below the critical Ca, the droplet size was independent of Ca. In contrast, above the critical Ca, the droplet size increased sharply with increasing Ca. The effect of viscous force may become signiﬁcant in this Ca range. It is also worth noting that the size and size distribution of the monodisperse emulsions produced by MCE is not sensitive to the ﬂow rate of the dispersed phase below the critical value (Figure 7.4) [55]. Monodisperse emulsions can be stably produced by MCE when the continuous phase preferentially wets the surface of an MCE chip [51, 55, 74, 75]. The electrostatic interaction between the chip surface and emulsiﬁer molecules also critically affects droplet generation from MCs. Uniform droplets can be generated when the MCE chip used has a nonattractive interaction with emulsiﬁer molecules. It is important to keep the charge of the chip surface and emulsiﬁer used in mind during MCE as well as ME.

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Figure 7.4 Effect of the dispersed-phase flow rate on the size and size distribution of the produced O/W emulsions [55]. Refined soybean oil was used as the dispersed phase, and Milli-Q water containing 1 wt% sodium dodecyl surface (SDS), as the continuous phase.

7.3 Production of Structured Systems for Food Applications 7.3.1 O/W Emulsions 7.3.1.1 Membrane Emulsification O/W emulsion is important on the formulation of many food products. Food emulsions normally require droplet size in the range of 0.1 and 30 mm, mayonnaise, salad dressing, cream liqueurs, and ice cream, milk and dairy drinks. An overview on the production of single and multiple emulsions, gel microbeads, solid lipid microparticles, protein microspheres by ME was reported by Vladislavljevic and Williams [7, 76]). You et al. [77] have produced gel particles as calcium alginate using ME, where a microporous glass membrane of 2.9 mm average pore diameter was

7.3 Production of Structured Systems for Food Applications

used to produce calcium alginate microspheres with 4 mm mean diameter. Another example is Liu et al.s [78] work, where they used metallic membrane with pore diameter raging from 2.9 to 5.2 mm to produce uniform calcium alginate spherical beads with mean diameter of 50 mm. Both groups observed that transmembrane pressure is one of the most important process parameters to determine ﬁnal particle size and its size distribution. Solid lipid particles were produced by ME, as shown by Doria et al. [79]. Mean particles size from 50 to 750 nm could be reached at disperse phase ﬂow rates up to 0.84 m3 m2 h1. Fouling limited the production rate when membranes with small pore diameter (0.2 and 0.4 mm) were used. As a food formulation Gijsbertsen-Abrahamse et al. [37] have computationally simulated a culinary cream containing 30% fat using three different type of membranes (SPG, ceramic and microsieve) with 0.2 mm pore diameter. A microsieve membrane provided the best performance for an industrial scale production due to its lowest porosity (e ¼ 0.01) compared to SPG (e ¼ 0.6) and ceramic (e ¼ 0.35) membranes. 7.3.1.2 Microchannel Emulsification Much of the MCE literature has discussed the production of O/W emulsions consisting of food-grade substances. The surfaces of MC arrays as well as the membrane surfaces must remain sufﬁciently hydrophilic during MCE. Prior to ﬁrst usage, MC emulsiﬁcation chips made of single-crystal silicon are subjected to plasma oxidation in order to grow a hydrophilic silicon dioxide layer on the surface of the MC arrays [55]. Vegetable oils (reﬁned soybean oil and high oleic sunﬂower oil) and medium-chain triglyceride (MCT) have been used as the dispersed phase for producing monodisperse O/W emulsions by MCE [51, 80]. Tan et al. [81] also demonstrated that monodisperse O/W emulsions are produced when reﬁned palm olein is used after removing monoglycerides and diglycerides. Thus, it is necessary to keep in mind that the hydrophobicity of the dispersed oil phase is a critical parameter affecting the generation of oil droplets in MCE. Since food-grade oils are generally viscous liquids at room temperature, the ratio of the dispersed-phase viscosity to the continuous phase viscosity is usually high unless the continuous water phase contains a considerable amount of thickeners. This high viscosity ratio has the merit that droplet generation is not sensitive to operating conditions. The ﬁrst generation of MCE chips, called grooved MC array chips, have a throughput of vegetable oil droplets of less than 1 L m2 h1. Kobayashi et al. [54] developed straight-through MC array chips as high-throughput MCE chips, realizing the generation of uniform vegetable oil droplets at a high dispersed-phase ﬂux of up to 65 L m2 h1. In MCE, monodisperse O/W emulsions can also be produced using chemical oils (e.g., alkane oils and silicone oils) with a wide viscosity range of 1 mPa s to 103 mPa s and a continuous water phase with a viscosity of 1 mPa s [82, 83]. Droplet production per MCE chip tends to increase as the dispersed-phase viscosity decreases, indicating that MCE at an elevated temperature can increase the production of vegetable-oil droplets. The effect of food-grade emulsiﬁers on MCE has been reported several times in the literature. Several nonionic emulsiﬁers (TweenÒ 20, TweenÒ 80, pentaglycerin

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monolaurate, and sucrose monolaurate) have been demonstrated to be appropriate for stably generating uniform droplets by MCE [75, 84, 85]. Although uniform vegetable oil droplets were generated from an MC array in the absence of emulsiﬁers, the generated droplets were unstable, as was expected [84]. The results reported in the literature suggest that hydrophilic nonionic emulsiﬁers with a hydrophile– lipophile balance (HLB) exceeding 10 must be used in order to stably produce monodisperse O/W emulsions by MCE. Tong et al. [86] investigated the production of O/W emulsions stabilized by phospholipids, demonstrating that uniform oil droplets were generated using a continuous water phase containing anionic lyzophosphatidylcholine (LPC). Interestingly, droplet generation was made more stable by using lecithin in the dispersed phase and LPC in the continuous phase. The effect of proteins as an emulsiﬁer on MCE was also investigated by Saito et al. [87]. Dropletgeneration behavior (Figure 7.5) was found to be highly relevant to protein solution properties, such as the isoectronic point (pI), contact angle, and interfacial tension. When the pH of the continuous water phase was close to 7, MCE generated uniform vegetable oil droplets stabilized by bovine serum albumin (BSA), b-lactoglobulin, soybean ﬂour, or whey protein, which have a low pI, a high contact angle of an oil droplet, and/or low interfacial tension. In contrast, no droplets were generated for a continuous water phase containing lysozyme (pI: 10.5–11.0) or egg-white protein. It is also important to control the pH of the continuous water phase during MCE, since

Figure 7.5 Generation of soybean oil-in-water emulsion droplets stabilized by proteins from MCs [87]. (a) and (b) Generation of uniform droplets. (c) Unstable generation of nonuniform droplets. (d) Wetting of the dispersed phase on the chip surface.

7.3 Production of Structured Systems for Food Applications

the charge of the protein molecules reverses close to pI. As described in Section 7.2, the charge of an emulsiﬁer greatly affects its interaction with the negatively charged surface of an MC array as well as the droplet generation. Uniform droplets stabilized by negatively charged BSA molecules were generated at pH values over pI at 4.7–4.8, whereas the dispersed phase covered by positively charged BSA molecules wetted on the chip surface at pH values below pI. Food-grade O/W emulsions produced by MCE have been applied to produce monodisperse microparticles and microcapsules. Sugiura et al. [88] obtained dispersions of monodisperse solid lipid microparticles by cooling uniform droplets of melted oils (tripalmitin and hydrogenated ﬁsh oil) generated using MCE. Kobayashi et al. [89] produced dispersions of tripalmitin microspheres by MCE and subsequent solvent evaporation. The MCE in this work produced monodisperse O/W emulsions consisting of hexane, which was chosen as a solvent. Although dichloromethane is commonly used for solvent evaporation, this solvent is not available for food applications. The hexane in the oil droplets successfully transferred to the continuous water phase during solvent evaporation at atmospheric pressure and room temperature, considerably reducing the droplet size (Figure 7.6). Nakagawa et al. [90] produced dispersions of monodisperse gelatin/acacia complex coacervate microcapsules by MCE and subsequent coacervation. The single-core microcapsules were prepared using uniform vegetable oil droplets stabilized by appropriate types of gelatin. Chuah et al. [91] formulated monodisperse O/W emulsions stabilized by a layer of an electrolyte complex of negatively charged modiﬁed lecithin and positively charged chitosan (Figure 7.7). Uniform vegetable oil droplets stabilized by modiﬁed lecithin were initially generated by MCE. Adding a sufﬁcient amount of chitosan to the preceding O/W emulsions yielded positively charged oil droplets, with higher stability against heating (particularly at 70–90 C) and long-term storage at pH 3 than oil droplets stabilized solely by modiﬁed lecithin. The above-mentioned monodisperse microparticles and microcapsules produced by this process are promising for food applications, although their production scale is currently less than 1 g h1. Their throughput must be scaled up for practical-scale production.

Figure 7.6 Size reduction of uniform oil droplets by solvent evaporation [89]. (a) Optical micrograph of uniform hexane oil droplets containing tripalmitin produced by MCE. (b) Optical micrograph of uniform tripalmitin microspheres after solvent evaporation.

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Figure 7.7 (a) Schematic drawing of the formation process of electrostatic complex on the droplet surface. (b) Optical micrograph of uniform soybean oil droplets stabilized by a thin layer of electrostatic complex of modified lecithin and chitosan [91].

7.3.2 W/O Emulsions 7.3.2.1 Membrane Emulsification Most of the ME/MC literature concerns oil-in-water (O/W) emulsions, however, the production of water-in-oil (W/O) emulsions has also been discussed [16]. The basic principles set out in the previous chapters apply to both cases, but in practical terms the preparation of W/O emulsions with ME (both XME and PME) differs in two important aspects from the O/W case. First, a hydrophobic membrane surface has to be provided and maintained. Secondly the viscosity ratio of dispersed and continuous phase can be quite different. The latter may not be very signiﬁcant when low-viscosity hydrocarbons and water are involved (viscosity ratio not far from unity), but for food emulsions based on vegetable oil (50–60 mPa s at room temperature) the viscosity ratio can differ by up to two orders of magnitude between an O/W and W/O emulsion made from the same materials, unless the water phase is thickened considerably.

7.3 Production of Structured Systems for Food Applications

Providing a hydrophobic surface can be done by modiﬁcation of the usual hydrophilic membranes like Shirasu-porous-glass (SPG) membranes. The papers by Cheng et al. [92, 93] provide recent examples of this approach, in which a silane coupler or a silicone resin was used to render the surface of SPG membranes hydrophobic. Also, silicon nitride microsieves and perforated steel plates have been made hydrophobic via chemical surface treatment [94, 95]. Katoh et al. [15] have simply soaked SPG membranes thoroughly in the oil phase prior to the ME preparation of W/O emulsions. Sotoyama et al. [96] have also used this approach, and suggested that the added emulsiﬁer in the oil (in their case polyglycerin polyricinolate, PGPR) adsorbs to the silanol groups on the glass surface creating a hydrophobic base. This soaking procedure is potentially an attractive option for food applications as there is no risk of gradual wear of adsorbed chemicals from the membrane. However, more work will be needed to determine how strong the effect is for a given oil þ emulsiﬁer and how long the effect persists. Another option for getting the right wetting behavior is to use membranes made of a hydrophobic material like polypropylene [97] or polytetraﬂuoroethylene [98]. The latter authors investigated both kerosene and corn oil (with emulsiﬁers) as the oil phase, and found that preparation of a corn-oil-based W/O emulsion was not possible at moderate transmembrane pressure. This was attributed to the wetting behavior of the PTFE, which appeared to be nonwetting for the corn oil. It is noteworthy that Vladislavljevic et al. [97] and Yamazaki et al. [98] have also used presoaking with the continuous oil phase, and that the former authors found a signiﬁcant effect on droplet size and dispersed phase ﬂux at given transmembrane pressure. 7.3.2.2 Microchannel Emulsification The production characteristics of W/O emulsions using MCE have been reported several times in the literature. A prerequisite for producing monodisperse O/W emulsions by MCE is to keep the surface of MC arrays hydrophobic, similar to ME. Hydrophobic treatment of silicon MCE chips is conducted by modifying their hydrophilic surface using a silane-coupler reagent [51, 99]. Liu et al. [100] and Kobayashi et al. [101] have developed MCE chips made of a naturally hydrophobic polymer (poly(methyl methacrylate), PMMA). PMMA grooved MC arrays were fabricated by injection molding [100], and PMMA straight-through MC arrays were fabricated as part of the (Lithographie, Galvanoformung, Abformung (LIGA) process [102]. MCE chips made of PMMA as well as membranes made of hydrophobic materials do not require any chemical surface modiﬁcation, which is advantageous for food applications. However, polymeric MCE chips are not strong against organic solvents frequently used as the continuous oil phase. In MCE, W/O emulsions have generally been produced using alkane oils with low and medium carbon numbers as the continuous phase due to their low viscosity. Only two studies have reported the production of water-in-triglyceride emulsions using MCE [99, 103]. Uniform water droplets could be generated in the absence of thickeners in triglyceride oils (MCT oil, soybean oil, or triolein oil) as the continuous phase via MC arrays (Figure 7.8). As mentioned earlier, the viscosity ratio of the dispersed water phase to the continuous triglyceride-oil phase decreases by two or

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Figure 7.8 Production of monodisperse water-in-triglyceride oil emulsions stabilized by a hydrophobic emulsifier using MCE. The oils used were MCT oil (a) and refined soybean oil (b). The osmotic pressure of the dispersed phase was 4.2 MPa [99].

three orders of magnitude compared to O/W emulsions made of the same two liquids. Optical microscopy during MCE suggests that droplet generation using water-in-triglyceride oil systems is less stable and more sensitive than water-in-alkane oil systems, a ﬁnding that could be attributable to the very low viscosity ratio of the former systems. Moreover, the use of the viscous continuous phase results in a quite low droplet-generation rate per MC, which can be somewhat increased by operating the MCE at an elevated temperature. For the effect of the dispersed water phase, its osmotic pressure is a critical parameter affecting droplet-generation behavior. Kobayashi et al. [99] clearly demonstrated that monodisperse W/O emulsions were stably generated at osmotic pressures above a critical value and that nonuniform water droplets are unstably generated below the critical osmotic pressure (Figure 7.9). At low osmotic pressures, an aggregated layer driven by spontaneous emulsiﬁcation was formed around the expanding dispersed phase and the generated droplets (Figure 7.9), which may prevent smooth movement of the water/oil interface inside an MC array. Food-grade nonionic emulsiﬁers have been primarily used to produce W/O emulsions in MCE. Sorbitan fatty acid esters were successfully used to produce monodisperse W/O emulsions [74]. Sugiura et al. [103] screened polyglycerin fatty acid esters and polyglycerin condensed ricinoleic acid esters suitable for producing monodisperse W/O emulsions. Polyglycerin condensed ricinoleic esters with a very low HLB value (<1) were found to be particularly suitable for stably generating uniform water droplets. In contrast, the use of soybean and egg-yolk lecithins resulted in unstable generation of water droplets and their immediate coalescence. Until now, monodisperse gel microbeads and giant vesicles have been obtained using uniform water droplets produced by MCE as templates. Kawakatsu et al. [104] produced monodisperse albumin gel microbeads by denaturing droplets of albumin aqueous solution dispersed in a continuous oil phase. Iwamoto et al. [105] obtained monodisperse gelatin gel microbeads by cooling droplets of gelatin aqueous solution produced by MCE at an elevated temperature. The particle size hardly changed during gelation of the droplets containing albumin, whereas a signiﬁcant decrease in

7.3 Production of Structured Systems for Food Applications

Figure 7.9 (a) and (b) Effect of the osmotic pressureofthe dispersed phaseonthe generation of W/O emulsion droplets from MCs. (a) Generation of uniform aqueous droplets at an osmotic pressure of 4.2 MPa. (b) Unstable

generation of nonuniform Milli-Q water droplets. (c) and (d) Formation of aggregates around the generated Milli-Q water droplets. Optical micrographs of the resultant water droplets just after generation (c), after 20 min (d).

particle size occurred during gelation of the droplets containing gelatin. Monodisperse gel microbeads are considered to be promising microcarriers for functional food ingredients; however, more work has to be done to precisely control their particle size and to produce monodisperse gel microbeads encapsulating functional food ingredients. Monodisperse giant vesicles have also been obtained by the lipid-coated ice droplet hydration method using aqueous droplets dispersed in the continuous phase of a hexane solution generated by MCE (Figure 7.10) [106, 107]. The monodisperse giant vesicles consist of food-grade substances, but phosphatidylcholine used as an emulsiﬁer is very expensive; therefore, they would be promising for pharmaceutical applications. 7.3.3 W/O/W Emulsions 7.3.3.1 Membrane Emulsification The ﬁrst work on the production of multiple emulsions was published in 1923 [108]. Multiple emulsions are complex structures with special properties as carrier systems,

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Figure 7.10 (a) Optical micrograph of W/O emulsion droplets generated by MCE. (b) Images of a giant vesicle observed by bright-field light microscopy (left) and by fluorescence microscopy (right) [107].

and have been recently used in the manufacturing of low energy density food products. Muschiolik et al. [49, 109] have reported the production of multiple emulsions (W/O/W) by cross-ﬂow and rotating membrane emulsiﬁcation, single and double T-junction microchannel, and glass capillary (coaxial jet). More recent developments on the production of double emulsions using microﬂuidic devices, including membrane and microchannel emulsiﬁcation were published by Vladislavljevic and Williams [76] and others as cited in Section 7.3.2.2. Stability of multiple emulsions can be inﬂuenced by different factors as Laplace and osmotic pressures between internal and external phases, interaction between emulsiﬁers (low and high HLB), also between thickener and high HLB emulsiﬁer, and viscosity of both phases [110, 111]. Membrane emulsiﬁcation is a suitable process to produce multiple emulsions as process conditions are favorable due to low shear rates during processing. Their physical stability against Ostwald ripening, consequently long-term shelf life, depends on the balance between Laplace and osmotic pressures as previously mentioned. Such balance can be reached by adding salt to W1 (inner aqueous phase). A viscosity ratio of 1 between W2 (outer aqueous phase) and W1/O is preferable for the production of successful multiple emulsions. Thickeners, such as guar gum, xanthan gum, gelatin, maltodextrin (DE > 10), hydroxyethylcellulose. For the production of W1/O (dispersed phase), PGPG (polyglycerol ester of ricinoleic, HLB ¼ 4) and modiﬁed lecithin have been used for food applications, and nonfood grade cetyl dimethicone copolyol, PEG-30 dipolyhydroxystearate (block copolymer). W/O/W emulsions are promising structured systems for applications in the food industry for low-fat food formulations [112], and also delivery systems of (bio)active molecules. Kanouni et al. [110] has suggested the use of W/O/W emulsions on the formulation of sauce, mayonnaise, where a less oily taste may be reached. Another advantage is the formulation of low fat food products. Skin creams may also provide a different feeling after rubbing on skin. As an example, food-grade W/O/W emulsions were produce by dead-end membrane emulsiﬁcation with droplet size of 100 mm [4]. An SPG membrane was used for the production of narrow droplet-size distribution with the smallest span of about 0.28 at high ﬂow rate. Several passes allowed very narrow droplet-size distribution.

7.4 Encapsulation of Active Molecules

W/O/Wemulsions containing whey protein isolate in the water internal droplets were produce using the same process. Gelation of internal water droplets by whey protein provided the smallest particle size and narrowest particle-size distribution [113]. 7.3.3.2 Microchannel Emulsification Production of W/O/W emulsions using MCE has been investigated in a few studies. The W/O/W emulsions were produced by two-step emulsiﬁcation processes. As a ﬁrst-step emulsiﬁcation, W/O emulsions was prepared by homogenization [114, 115] or microﬂuidization [116]. Homogenization yielded W/O emulsions with an average droplet size on the order of several micrometers to several tens of micrometers. In contrast, microﬂuidization enabled the preparation of ﬁne W/O emulsions with an average droplet size as small as 150 nm. In second-step emulsiﬁcation (MCE), oil droplets containing smaller water droplets were generated by injecting a W/O emulsion into a continuous external water phase through MCs. During MCE, the hydrophilic surface of the silicon MCE chip must be maintained in order to produce monodisperse W/O/W emulsions. Water-in-triglyceride oil emulsions (soybean oil, triolein, or MCT-oil) have been successfully used to produce monodisperse W/O/W emulsions by MCE [114–116]. The monodisperse W/O/W emulsions were stabilized by two food-grade emulsiﬁers: a hydrophobic emulsiﬁer dissolved in the medium oil phase and a hydrophilic emulsiﬁer dissolved in the external water phase. In addition, the osmotic pressure of the internal and external water phases must be appropriately controlled in order to obtain stable W/O/W emulsions as well as feed W/O emulsions. Kobayashi et al. [116] demonstrated that the volume fraction of ﬁne water droplets dispersed in uniform oil droplets can be controlled and increased up to 30% (Figure 7.11). It is necessary to mention that no leakage of the internal water droplets was observed during MCE. Only one study has reported the production of dispersions based on the produced monodisperse W/O/W emulsions. Kawakatsu et al. [114] obtained food-grade S/O/W emulsions by electrolyte-induced gelation of the internal phase of a pectin aqueous solution. Currently, one can ﬁnd a much greater number of studies that discuss the production of W/O/W emulsions using ME as the ﬁrst-step and/or second-step emulsiﬁcation (see Section 7.3.3.1). In particular, information about two-step emulsiﬁcation processes using SPG membranes would be directly applicable to the production of W/O/W emulsions using MCE.

7.4 Encapsulation of Active Molecules 7.4.1 Membrane Emulsification

Membrane and microchannel emulsiﬁcation are gently technologies to encapsulate sensitive compounds into single and multiple emulsions, as well as microcapsules [7], due to its low shear rate during processing. Multiple emulsions are also

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Figure 7.11 (a) Schematic drawing of the production of a W/O/W emulsion by MCE. Size (b) Optical micrograph of the generation of uniform soybean oil droplets containing aqueous droplets from MCs. (c) Optical

micrograph of the generated oil droplets containing many submicrometer aqueous droplets. The volume fraction of the internal water phase in the oil droplets was 30% [116].

potential matrixes to encapsulate active molecules for inumerous applications in food, cosmetic, and pharma industries [76, 117, 118]. They have been used in the encapsulation of compounds as drugs [119], vitamins, retinyl palmitate, carotenoids, polyphenols [120], ﬂavors, ions Mg2 þ [121], and antimicrobials [8, 122]. These carrier systems may protect sensitive compounds against chemical, and enzymatic degradations; and mask undesirable taste. Literature has shown a successful incorporation of microorganisms into microcapsules by membrane emulsiﬁcation. Zhou et al. [123] studied the encapsulation of bacterial cells into uniform-sized agarose microcapsules by membrane emulsiﬁcation. Cell growth could be observed after 14 days of incubation time and it showed that this mild process was able to preserve cell viability. In another research work, uniform droplets and microcapsules containing Lactobacillus casei were also produced by ME [124] for further application in dairy products. Monodisperse W/O chitosan emulsion as insulin carrier systems was prepared by membrane emulsiﬁcation and followed by cross-linking using tripolyphosphate (TPP) and glutaraldehyde for two steps of particle solidiﬁcation [125]. Uniform-sized microspheres were able to keep insulin activity and provide high encapsulation efﬁciency. Ribeiro et al. [42] investigated the encapsulation of astaxanthin, a carotenoid, in O/W emulsion by dead-end membrane emulsiﬁcation. For the production of smaller

7.5 Assessment of the Potential Benefits of Membrane Emulsification in Foods

droplet size and narrower droplet-size distribution, three passes through the membrane were used. A strong fouling could be observed due to the protein used as the emulsiﬁer. 7.4.2 Microchannel Emulsification

Droplet generation for MCE is a very mild process driven by spontaneous transformation of the dispersed phase that passes through the MCs [103]. The energy input for MCE is also very low (e.g., 103–104 J m3), indicating that temperature elevation during emulsiﬁcation can be neglected [103]. These features are attractive for preventing the degradation of shear- and heat-sensitive active molecules. Monodisperse O/W emulsions consisting of hydrophobic active molecules using MCE have been recently produced by Neves et al. [85, 126]. Neves et al. [85] ﬁrst discussed the generation of soybean oil droplets containing beta-carotene of gamma-oryzanol from MCs. Uniform oil droplets containing gamma-oryzanol were generated in a continuous phase containing a food-grade hydrophilic emulsiﬁer at room temperature. For beta-carotene, MCE was conducted at an elevated temperature to prevent recrystallization of beta-carotene dissolved in the soybean oil, resulting in uniform oil droplets. Neves et al. [126] also generated droplets of reﬁned palm oil rich in betacarotene and ﬁsh oil droplets or a mixture of palm oil and ﬁsh oil rich in polyunsaturated fatty acids. Droplet production per MCE chip was almost independent of the concentration of the active molecules [126]. Sugiura et al. [106, 115] discussed the entrapment yield of model ﬂuorescent molecules (calcein) in a W/O/W emulsion and giant vesicles obtained using MCE. Hydrophilic calcein was added in the internal water phase before producing the W/O emulsions. The entrapment yield in the W/O/W emulsion was very high (91%), which is considered to be attributable to the very mild droplet-generation process via MC arrays [113]. Giant vesicles obtained by Sugiura et al. [106] had the highest entrapment yield of approximately 35%, comparable to the reverse-phase evaporation method, and was signiﬁcantly higher than most other giant-vesicle formation processes. In this case, 67% of the calcein leaked out from the internal water phase, mainly during the hydration step. A further modiﬁcation of the hydration process must be undertaken to achieve higher entrapment yields.

7.5 Assessment of the Potential Benefits of Membrane Emulsification in Foods

The potential beneﬁts of ME have been discussed by many authors, for example, Joscelyne and Tr€agård [3], Charcosset et al. [34] and quite recently Charcosset [9]. In fact, low shear, low power input and narrow droplet-size distribution (compared to conventional emulsiﬁcation) are mentioned as beneﬁts in the introduction of all papers on ME, but this is rarely discussed in more detail from a food-industry point of view. In this section we will try to ﬁll this gap to some extent by taking a food product

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developer perspective and assess what mild processing and a narrow droplet-size distribution (DSD) can actually be expected to contribute to the consumer-perceived properties of a food product. In our view this is less evident than most academic ME literature suggests, and realizing this should help to focus the efforts on the industrialization of ME for foods applications. 7.5.1 DSD and Product Stability

A narrow droplet-size distribution is frequently claimed to enhance product stability during shelf life, which may be divided into physical, chemical and microbiological stability. 7.5.1.1 Physical Stability Physical stability typically refers to two aspects: (1) changes in the DSD via coalescence or Ostwald ripening, and (2) creaming or settling of the droplet phase. In principle, the width of the DSD can indeed affect these processes. The difference in Laplace pressure between droplets of different size is the driving force for Ostwald ripening, so this process will be slow if the DSD is narrow. Furthermore, a distribution in droplet size implies a distribution in creaming/settling velocity. This promotes the occurrence of droplet collisions, which may enhance coalescence if the droplets are not well stabilized. Also, the overall creaming/settling rate can be enhanced, because larger droplets tend to drag smaller ones along in their slipstream [127]. While this argumentation is valid in principle, its practical importance depends on the food emulsion considered. In many food products creaming/settling is prevented anyway because the continuous phase is structured by fat crystals or gelling agents (e.g., margarine, dressings), because the volume fraction is so high that the closest packing is obtained (mayonnaise), or because the droplet phase is clustered into a space-ﬁlling network (certain creams). In fully liquid products that require long-term stability, the droplet size as such can be so small that Brownian motion counteracts creaming (cream liqueur). In other semiliquid emulsions the phase separation is simply accepted. In those cases the droplets are well stabilized against coalescence and the consumer is requested to shake the bottle before use. For most oil-in-water (O/W) products, the timescale for Ostwald ripening exceeds the product shelf life. This is due to a combination of low solubility of triglyceride oils in water and mass transfer limitations presented by protein layers at the interface. In semisolid W/O emulsions like margarines a fat crystal shell around the droplets provides a mass transfer limitation as well as a mechanical restriction on droplet size changes. In more liquid-like W/O emulsions (e.g., pourable margarine) an osmotic stabilization against Ostwald ripening can be provided by salt in the water phase [128], which is often there for taste reasons anyway. Summarizing, practical cases where a narrow DSD could help to solve an urgent problem with the physical stability of a food emulsion are not readily apparent.

7.5 Assessment of the Potential Benefits of Membrane Emulsification in Foods

7.5.1.2 Chemical Stability Lipid oxidation and the consequent production of off-ﬂavors is a general problem in products based on triglyceride oils. For O/W emulsions, the DSD at given volume fraction can play a role here in principle, as it determines the interfacial area between the oil and the water. In a review of lipid oxidation in O/W emulsions, McClements and Decker [129] pointed out that only a limited number of studies have been done into the effect of droplet size on oxidation. Some indeed corroborate the expectation that oxidation increases with decreasing droplet size at given volume fraction, due to the increase in interfacial area. One study found no effect. McClements and Decker [129] suggested that the presence of catalytic species with a preference for the interface could explain this observation. If all available catalyst species reside at the interface anyway, the interfacial area becomes unimportant. In any case reducing the total interfacial area by narrowing the DSD has only limited potential. Assuming a lognormal size distribution and typical parameters for conventionally produced emulsions we have estimated that a perfectly monodisperse emulsion of the same volume fraction would only have a 30% smaller total interface. 7.5.1.3 Microbiological Micro-organisms can grow in the water phase of the emulsion, which implies that droplet size is of primary important in W/O emulsions. Water droplets can be made small enough to suppress the growth of micro-organisms due to insufﬁcient amount of nutrient per drop and to space limitation within a small drop [130]. Obviously, the maximum droplet size is the key parameter here. The width of the DSD does not play a direct role. 7.5.2 DSD and Product Rheology

The rheology of dispersions and emulsions has been the subject of many textbooks and articles, and a full review is well beyond the scope of the present discussion. Rather, we want to focus on a limited number of references in which the role of polydispersity of the dispersed phase is considered. The most relevant rheological parameters for food-type emulsions are: . . .

viscosity; the linear viscoelasticity parameters G0 and G00 ; the yield stress.

All four parameters depend on the DSD, although in many food emulsions a signiﬁcant (if not dominant) contribution also comes from structure in the continuous phase. This will not be considered explicitly here. Often, relations that were originally derived for dispersions of solid particles are used. This is a good approximation when (1) the Laplace pressure is high enough compared to the applied hydrodynamic stress to prevent signiﬁcant droplet deformation, and (2) the droplet interface behaves quasirigidly with respect to tangential hydrodynamic stress, due to the presence of surface-active molecules or an interfacial

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ﬁlm/skin [131]. However, in practice these relations are also used as a ﬁrst estimate beyond the range of their strict applicability, because manageable models that take into account droplet deformability as well as interfacial rheology are not readily available. 7.5.2.1 Yield Stress In many food emulsions that possess a yield stress, this is due to structure in the continuous phase, for example, the fat–crystal network in margarines. An appreciable yield stress due to the dispersed phase is only observed in concentrated emulsions, in which the droplets are closely packed. Mayonnaise (O/W emulsion with 80% oil) is a well-known example. This raises the question at which volume fraction an emulsion becomes close-packed. For monodisperse solid spheres the maximum packing fraction depends on the type of packing, varying from 0.63 for random packing to 0.74 for a face-centeredcubic (FCC) crystalline lattice. However, both computer simulations and experiments have indicated that the close-packing value for dispersions of monodisperse spheres rarely exceeds the random-packing value, to be denoted henceforth as wRCP [131– 135]. Also for w > wRCP the structure often remains disordered, unless speciﬁc measures are taken to make it more regular (e.g., by application of a well-deﬁned ﬂow). Mason et al. [131], for instance, have demonstrated via light scattering that the quasimonodisperse emulsions in their experiments were all disordered on a macroscopic length scale, even at volume fractions close to unity. The consequences of the disordered microstructure of concentrated emulsions for their rheological properties have been discussed in detail by Mason et al. [131, 132]. When a stress is applied to a structure with ordered packing, yielding implies a global topological rearrangement, that is, planes of droplets move in unison. According to computer simulations cited by Mason et al. [131], this occurs at strains of the order 0.6 for a 3D ordered lattice. By contrast, a disordered structure allows yielding via local rearrangements of droplets or groups of droplets. This can take place already at much lower strain. Hebraud et al. [133] have given an elegant experimental conﬁrmation of this view by probing local rearrangements during the yielding of disordered emulsions via diffusing-wave spectroscopy. The agreement of their calculations and experimental results with the data of Mason et al. [131] is also quantitatively good. Mason et al. [131] already conjectured that it is the disordered microstructure of real emulsions rather than their polydispersity that explains the discrepancy between experimental data for polydisperse emulsions and theoretical results for particle packings of monodisperse emulsions. This issue has been considered in more detail by Saint-Jalmes and Durian [135], in a study of polydisperse foams (which are very similar in rheological behavior to concentrated emulsions). These authors found quantitative agreement with the correlation of Mason et al. [131], despite the polydispersity of their foams. Also, other rheological parameters were quite similar, which lead Saint-Jalmes and Durian [135] to the conclusion that polydispersity does not play an important role in concentrated systems, as long as it stays moderate and the DSD is unimodal.

7.5 Assessment of the Potential Benefits of Membrane Emulsification in Foods

7.5.2.2 Elastic Modulus Mason et al. [132] considered the elasticity of concentrated monodispersed emulsions, and found that the elastic shear modulus G0 scales as w(w wRCP)s/R, where s is the interfacial tension and R the droplet radius. As for the yield stress, Saint-Jalmes and Durian [135] demonstrated that this correlation also holds for polydisperse foams. Again, their conclusion was that the disordered structure of the foams makes polydispersity of minor importance, and this reasoning can be extended to emulsions. 7.5.2.3 Viscosity The above rheological properties characterize the resistance of the system to stresses that tend to induce ﬂow. For a ﬂowing emulsion the apparent emulsion viscosity is the most important parameter. The viscosity of a dispersion can be phenomenologically related to the DSD via the well-known Krieger–Dougherty equation, which links the viscosity to the actual and maximum volume fraction of dispersed phase [134, 136]: mREL ¼

m w ½mwM ¼ 1 mC wM

Here, mC is the viscosity of the continuous phase, [m] is the so-called intrinsic viscosity, and wM is the maximum packing fraction of the droplets. In fact, Barnes [136] suggested that the exponent [m]wM is often close to 2. The relative emulsion viscosity is thus sensitive to the precise value of the maximum packing fraction, which depends on the DSD. Figure 7.12 shows the typical range encountered in food emulsions. The effect becomes signiﬁcant for volume fractions above about 0.4. A successful ﬁt of the viscosity to the KD equation, using the maximum packing fraction as a ﬁt parameter, does as such not highlight the physical background of the processes involved. One might ask how particles at volume fractions below the closest packing would know what their wM would be upon increasing the volume fraction. Actually the mechanism of viscosity increase with volume fraction is based on hydrodynamic interactions between the particles, which have a size dependence and thus give a relation to the DSD. Apparently, this can be adequately captured phenomenologically by choosing the wM that corresponds to the DSD at hand. 7.5.2.4 Formation of Flocculated Networks The effect of DSD width on aggregation/ﬂocculation has been considered by Bushell and Amal [137] in computer simulations based on diffusion-limited cluster aggregation. They found that the fractal structure and the form of the function that describes the gross shape of the aggregates is unaffected by details of the primary particle-size distribution. Bushell and Amal [137] claimed that their results are consistent with other literature on the effect of polydispersity on aggregation (see their paper for references).

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0.63

0.7

0.75

0.8

80 70

Relative viscosity [-]

60 50 40 30 20 10 0 0

0.2

0.4

0.6

0.8

volume fraction [-] Figure 7.12 Krieger–Dougerthy (KD) relation for different values of the maximum packing fraction.

7.5.3 Product Properties Related to Low-Shear Processing 7.5.3.1 Shear Damage to Ingredients ME is often claimed to give less deterioration of delicate ingredients. However, the standard macromolecular ingredients like proteins, enzymes, and polysaccharides are quite stable against shear damage for typical conditions in conventional industrial emulsiﬁcation, except when homogenization is done such that cavitation is present. Homogenization pressures required to affect such ingredients are much higher, and are then deliberately used to change the functionality of the macromulecules [138–141]. Besides the standard ingredients one can think of additives for functional foods, which can be dissolved molecules or particulates. The former quite probably have at least the shear stability of macromolecules. The latter might be encapsulated goodies or microbiological cells (living cultures). It is difﬁcult to make general comments on the stability of encapsulates, given their variety and the dependence of their strength on product conditions like moisture level and pH. Typically, encapsulates will not be broken by pure simple shear, since their apparent viscosity ratio will be large and they will exhibit solid-body rotation rather than deformation and break-up. Added cultures usually are so-called gram-positive bacteria, which have a strong cell wall. Indeed the very high pressure homogenizers mentioned above are typically needed for the disruption of such cells [138].

7.5 Assessment of the Potential Benefits of Membrane Emulsification in Foods

7.5.3.2 Effect on Product Structure It has been demonstrated in the literature that XME is a suitable method to produce double emulsions (oil-in-water-in oil, O/W/O, or water-in-oil-in-water, W/ O/W) [142]. XME is then used to disperse the primary emulsion ﬁnely into the outer phase. The success of XME in this application is directly related to the mild processing conditions. In conventional emulsiﬁcation, the requirement to have a rather small size of the droplets of the outer emulsion (W/O droplets in W/O/W) implies the use of high shear. Muguet et al. [143] have demonstrated that this implies an increased release of internal droplets. The Japanese company Morinaga has patented the use of XME for making duplex spreads in the early 1990s [110], although we are not aware that they actually have a product on the market. In many cases, a structure is building up in the emulsion during emulsiﬁcation. For instance, a fat–crystal network is starting to form in the oil phase of a margarine emulsion during cooling and emulsiﬁcation in scraped-surface heat exchangers. In certain creams, a network of aggregated droplets is formed. When biopolymer mixtures are present, structure formation via phase separation and/or gelation can occur. These structure-formation processes are all affected by shear, and as can be expected the high shear required to set the droplet size is not always desired from the structure formation point of view. Using a mild emulsiﬁcation method like XME might be beneﬁcial in this context, although the structure formation in the continuous phase may well cause problems to mix the droplets uniformly into it. 7.5.4 Summary

It is frequently claimed or suggested in the ME literature that having a narrow DSD provides signiﬁcant improvements in the properties of food products. The above considerations show that this is in fact far from obvious for many commercial food products. First, the properties of many products depend at least in part on the thickened and/or gelled continuous phase. Secondly, when focusing on the role of the DSD, we see that neither emulsion stability nor its rheology is signiﬁcantly affected by the width of the DSD, except possibly the apparent emulsion viscosity in some cases. The claimed beneﬁts of mild processing can be relevant for delicate ingredients and (micro)structures like encapsulated nutrients and ﬂavors, but the standard ingredients like proteins and polysaccharides survive the shear in conventional processing quite well. An interesting area for ME is that of multipleemulsion formation, as has also been pointed out by Charcosset [9]. Energy saving in emulsiﬁcation has been demonstrated, but in foods its signiﬁcance in the overall cost breakdown may often not be large enough to justify, on its own, a switch from conventional emulsiﬁcation to ME. Moreover, it has to be noted that total energy expenditure needs to be considered, that is, it may be that large-scale ME lines require more frequent cleaning than conventional emulsiﬁcation equipment.

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7.6 Conclusions

ME and MCE are low energy input processes and have been successfully applied for the precision manufacture of particulate systems, however, efforts are required to develop further these worthwhile technologies for large-scale production. One of the biggest challenges is the development of novel surface properties of the membranes, needed to control their surface energy and avoid changes in wetting properties over time. Another important aspect concerns the reduction of membrane fouling during processing. It would be a big disadvantage if preﬁltration of the droplet phase (in XME/MCE) or even of both phases (in PME) is needed, or if cleaning of the line needs to be done much more frequently than for conventional emulsiﬁcation equipment. Increase in volume production could make these technologies competitive to conventional mechanical emulsiﬁcation processes, allowing their process intensiﬁcation and a sustainable production. In several industries (e.g., pharmaceutical and ﬁne chemicals) a narrow dropletsize distribution can be advantageous, as discussed in many review papers. However, in many food products the droplet-size distribution does not play a dominant role, and the size distributions that can be reached with conventional equipment are adequate. ME and MCE have potential for energy saving, but this can only be assessed fully if also changes in cleaning procedures are taken into account. If the latter does not add much, a quite signiﬁcant energy saving for the emulsiﬁcation process (close to an order of magnitude) seems achievable. Relative to the total energy expenditure in food manufacturing (which also includes energy-intensive steps like pasteurization) this may not be large, but in absolute terms it will contribute to a reduction of the carbon footprint. Both ME and MCE processes have potential for the production of duplex emulsions. Up to now this has only been demonstrated at quite small scale, but a successful scale-up could lead to a range of novel food products.

References 1 Charpentier, J.-C. (2007) Modern

chemical engineering in the framework of globalization, sustainability, and technical innovation. Ind. Eng. Chem. Res, 46, 3465–3485. 2 Schr€ oder, V. (1999) Herstellen von Öl-inWasser-Emulsionen mit mikropor€ osen Membranen, PhD Thesis, Universit€at Karlsruhe, Shaker Verlag. 3 Joscelyne, S.M. and Tr€ agård, G. (2000) Membrane emulsiﬁcation – a literature review. J. Membrane Sci., 169, 107–117. 4 Vladisavljevi c, G.T., Shimizu, M., and Nakashima, T. (2004) Preparation of

monodisperse multiple emulsions at high production rates by multi-stage premix membrane emulsiﬁcation. J. Membrane Sci., 244, 97–106. 5 van der Graaf, S., Schro€ en, C.G.P.H., and Boom, R.M. (2005) Preparation of double emulsions by membrane emulsiﬁcation – a review. J. Membrane Sci., 251, 7–15. 6 Schubert, H. (2005) Emulgiertechnik – Grundlagen, Verfahren und Anwendungen, Behrs Verlag, Hamburg. 7 Vladisavljevi c, G.T. and Williams, R.A. (2005) Recent developments in manufacturing emulsions and particles

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8 Membrane Contactors in Integrated Processes for Fruit-Juice Processing Alfredo Cassano and Enrico Drioli 8.1 Introduction

The overall market for fruit juices has grown substantially in recent years probably due to public perception of juices as a healthy natural source of nutrients and increased public interest in health issues. Indeed, epidemiological studies have established a positive association between the intake of fruit and vegetables and a reduced rate of heart diseases mortality, common cancers and other degenerative diseases [1]. This protective role may be related to phytochemicals acting as antioxidants, freeradical scavengers and saviors of the cell. These biologically active compounds may act independently or in combination as anticancer compounds by different mechanisms and are better absorbed from juices than from plant tissues. However, it is well known that naturally occurring antioxidants could be signiﬁcantly lost as a consequence of processing and storage. Processing operations, such as peeling, cutting and slicing and thermal treatments induce rapid depletion in natural antioxidants in food [2]. Therefore, in order to preserve the quality of fruit juices the food industry has focused on the development of new processing techniques for minimally processed fruit and vegetable products. Traditional membrane processes such as enzyme membrane reactors (EMRs), microﬁltration (MF), ultraﬁltration (UF), nanoﬁltration (NF) and reverse osmosis (RO) are today key processes in the food industry for concentration, fractionation and puriﬁcation of liquid foods. Their intrinsic properties (low operating temperature, no special chemicals required, no phase changes involved, easy scale-up and modularity, uncomplicated operation and possibility of automation) make them a valid alternative to traditional methods of liquid foods treatment. Additionally, potential energy savings derived by membrane processes application in the food and drink industry can be estimated as 50%, as reported by Eichhammer [3]. The introduction of these technologies in fruit-juice processing represents one of the technological answers to the problem of the production of juices with high quality, natural fresh taste and additive free. Juice clariﬁcation, stabilization, depectinization, fractionation and concentration are typical steps successfully realized by using

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EMRs, MF, UF, NF and RO. In particular, UF and MF represent a valid alternative to the use of traditional ﬁning agents (gelatin, bentonite and silica sol) and ﬁlter aids in fruit-juice clariﬁcation, color removal and stabilization [4, 5]. Basically the juice is treated after enzymatic pulping. EMRs are new approaches in which, by choosing an effective membrane conﬁguration, pulping and clariﬁcation can be realized in one step [6]. Fruit juices are usually concentrated in order to reduce storage, package and shipping costs [7]. In addition, concentrated fruit juices, because of their low water activity, have a higher stability than single-strength juices. The concentration of fruit juices is usually obtained by multistage vacuum evaporation; however, this process results in a loss of fresh juice ﬂavors, color degradation and a cooked taste, recognized as off-ﬂavors, due to thermal effects. Alternative techniques, such as freeze concentration systems (cryoconcentration), in which water is removed as ice rather than as vapor, allow preservation of the aroma compounds but they are characterized by high energy consumptions [8]. Besides, the achievable concentration (about 50 Brix) is lower than the values obtained in thermal evaporation (60–65 Brix). The concentration of fruit juices by RO has been of interest in the fruit-processing industry for about 30 years. The advantages of RO over conventional concentration techniques are in terms of low thermal damage of the product, reduction of energy consumption and lower capital investments [9] as the process is carried out at low temperatures and it does not involve phase change for water removal. Most studies concerning the concentration of fruit juices (including apple, pear, grapefruit, kiwi, pineapple, passion fruit, tomato juice, etc.) by RO have mainly focused on the effect of membrane type and operating conditions on the retention of juice components and permeate ﬂuxes [10–19]. However, the osmotic pressure of the juice increases rapidly with the increasing of the sugar concentration (100 and 200 bar for concentrations of 42 and 60 Brix, respectively). The concentration also determines an increasing viscosity. Both factors inﬂuence the RO process, so the ﬁnal concentration cannot be higher than 20%: otherwise the process is not convenient from an economical point of view. For these limitations, RO can be considered an advantageous technique as a preconcentration step [20]. The separation and concentration of polyphenolic compounds from apple juice by using 1- and 0.25-kDa molecular weight cut-off spiral-wound NF membranes has been reported by Saleh et al. [21]. The concentration of apple and pear juices by NF at low pressures (between 8 and 12 bar) has also been investigated [22]. Other membrane processes such as pervaporation (PV) and electrodialysis (ED) and gas separation (GS) have been studied in the fruit and vegetable sector. ED is a promising method for juice deacidiﬁcation able to preserve the organoleptic properties of the juice and to produce valuable by-products such as citric acid [23]. PV constitutes a promising alternative to traditional techniques, such as distillation and partial condensation, for aroma recovery from fruit juices [24]. Membrane contactors (MCs) represent innovative membrane-based operations that, due their potential advantages, are considered as new interesting perspectives for industrial and scientiﬁc applications. In the ﬁeld of fruit-juice processing the integration of MCs with conventional membrane operations emerges as

8.2 Membrane Contactors: Fundamentals

an interesting opportunity due to the synergistic effects that can be achieved. As a matter of fact this integration makes MCs very competitive against conventional energy-intensive techniques (i.e., distillation and evaporation) in terms of energy consumption, product recovery and improvement of quality. Moreover, industrial cycles can be redesigned according to a process-intensiﬁcation strategy that aims at minimizing environmental impact, increasing safety, improving remote control and automation, and reducing production costs and equipment size [25, 26]. In this chapter the main properties of MCs of interest in fruit-juice processing are described. Their potentialities within integrated membrane systems as well as their main drawbacks related to their further implementation at the industrial level will be also discussed.

8.2 Membrane Contactors: Fundamentals

Membrane contactors are systems in which the membrane acts as a barrier between two phases (gas/liquid or liquid/liquid) permitting mass transfer of the components without dispersion of one phase within another. Unlike traditional pressure-driven membrane processes, membrane contactors are not selective towards particular components and the separation is based on the principles of phase equilibrium. Basically, the two phases are kept in contact through a microporous membrane in correspondence of the pore mouths, where the interface is established, and the species are transferred from one phase to the other by simple diffusion through the membrane pores [27]. Membranes used in MCs can be both hydrophobic and hydrophilic. Polypropylene (PP), polyvinylidene diﬂuoride (PVDF), polytetraﬂuoroethylene (PTFE), polyethylene (PE) and perﬂuoropolymers (e.g., hyﬂon) are typical hydrophobic polymers used for these applications. They can be wetted by nonpolar solutions, while the polar phase cannot enter into the membrane pores; in order to avoid dispersion phenomena, the pressure of the polar phase has to be equal to or higher than the pressure of the wetting phase. Moreover, in order to prevent the penetration of the polar phase into the pores, and consequently, a loss of membrane hydrophobicity, the critical penetration pressure should not be exceeded. For a speciﬁc material the critical penetration pressure (DP) depends on the liquid surface tension, the pore radius and the contact angle, as reported in the Laplaces equation: DP ¼ 2c

cos q r

ð8:1Þ

where c is the surface tension of the liquid, q the contact angle between the liquid and the membrane, r the radius of the pore. According to this equation, membranes with large pore sizes guarantee lower values of the critical penetration pressure and, consequently, a maintenance of their hydrophobic character, which is very important for a good performance of the system, especially in long-term applications.

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When the membrane is hydrophilic, the nonpolar phase remains blocked at the pore mouth, while the polar phase wets the membrane pores. In this case the pressure of the nonpolar phase must be equal to or higher than the polar phase pressure to avoid dispersion between the phases. The interface is formed at the pore mouth of the nonpolar phase side and it is maintained if the critical penetration pressure is not exceeded [28]. Typical advantages of MCs over conventional technologies (such as strippers, scrubbers, distillation columns, evaporators, etc.) are in terms of high and constant speciﬁc interfacial area, use of plastic equipments, high modularity and compatibility, easy scale-up and control, independence of the ﬂuid phases in contact, possibility to operate at room temperature, no ﬂooding, loading and foaming. On the contrary, drawbacks are mainly related to the presence of an additional resistance offered by the membrane, membrane fouling, operative pressures depending on critical penetration ones, limited lifetime of the membranes, high replacement costs and channeling of ﬂuids [25, 29]. Figure 8.1 shows basic conﬁgurations of MCs that can be employed in fruit-juice processing; they include membrane distillation, osmotic distillation and liquid supported membranes. These processes can be integrated in the production lines

Hydrophobic membrane

evaporation (a)

condensation

vapour

feed

Tf > T p

permeate

Hydrophobic membrane

evaporation (b)

condensation

vapour

feed

awf > awp stripping solution liquid membrane

(c)

stirred aqueous feed phase

stagnant film

stirred aqueous strip phase

Figure 8.1 Schematic representation of membrane contactors: (a) membrane distillation; (b) osmotic distillation; (c) supported-liquid membrane.

8.3 Osmotic Distillation

together with conventional membrane operations in order to achieve advanced molecular separations overcoming existing limits of the traditional membrane processes (i.e., osmotic pressure limits in RO). Membrane distillation (MD) and osmotic distillation (OD) are typical processes for fruit-juice concentration. In these processes, the driving force for mass transfer is induced by a vapor-pressure difference across the membrane [30]. Supported-liquid membranes (SLMs) have been proposed for the extraction of organic acids from fruit juices. In SLMs the two sides of a microporous support are in contact with two aqueous phases: a feed phase and a strip phase, respectively. A carrier transports a speciﬁc substance through the micropores from the feed to the strip side. 8.3 Osmotic Distillation 8.3.1 Process Fundamentals

Osmotic distillation is a membrane-contactor technique also known as osmotic evaporation, membrane evaporation, isothermal membrane distillation or gas membrane extraction. Its main advantage lies in its ability to achieve high concentrations, working at low temperature and pressure, thus avoiding mechanical damage and thermal degradation of the solutes [31]. The OD process is based on the use of a macroporous hydrophobic membrane separating two circulating aqueous solutions of different solute concentration: a dilute solution on one side and a hypertonic salt solution on the other side. The hydrophobic nature of the membrane prevents penetration of the pores by aqueous solutions, creating a vapor/liquid interface at each entrance of the pores. The difference in solute concentration, and consequently in water activity between the two solutions, induces a vapor-pressure gradient at the vapor/liquid interfaces that constitutes the driving force of the water transport from the high vapor pressure phase to the low one [32]. The water transport through the membrane can be summarized in three steps: (1) evaporation of water at the dilute vapor/liquid interface; (2) diffusional or convective vapor transport through the membrane pore; (3) condensation of water vapor at the membrane/brine interface [33–36]. The water vapor pressures at the pore mouths are related to the temperature and activities prevailing in the liquids facing the membrane by: Pw1 ¼ Pw* aw1

ð8:2Þ

Pw2 ¼ Pw* aw2

ð8:3Þ

in which Pw* represents the vapor pressure of pure water and aw the water activity in the solutions. The driving force (DPw ¼ Pw1 Pw2) for water transport is sustained by the activity difference Daw ¼ aw1 aw2.

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The stripping solution, after its dilution, can be reconcentrated by evaporation and reused in the OD operation. Therefore, it should be thermally stable and also preferably nontoxic, noncorrosive and of low cost. Consequently, salts showing large increases in solubility with temperatures and low equivalent weights are preferred since they can be evaporated at high concentrations without risk of crystallization in the evaporator. A number of salts such as MgSO4, NaCl, CaCl2, MgCl2, KH2PO4, K2HPO4 is suitable. NaCl is characterized by a relatively low water solubility and a rather low temperature coefﬁcient of solubility, while CaCl2 is sensitive to precipitation in the presence of CO2; further, they are quite corrosive to ferrous alloys at elevated temperatures. Potassium salts of ortho- and pyrophosphoric acid offer several advantages, including low equivalent weight, high water solubility, steep positive temperature coefﬁcients of solubility and safe use in foods and pharmaceuticals. The equivalent weight of salts and their water solubilities as well increase in the order NaCl > CaCl2 > K2HPO4. Since the osmotic activity of a salt is considered as the ratio between its water solubility and its equivalent weight, K2HPO4 and CaCl2 can be considered as better osmotic agents than NaCl [34]. It is known that fruit and vegetable juices contain small concentrations of volatile aroma compounds, which are generally lost by thermal evaporation reducing remarkably their quality. In the OD process the concentration is performed at low temperatures and the vapor pressure of these volatile compounds is depressed, reducing the driving force for their transport across the membrane. Furthermore, the solubilities of these solutes in pure water are much higher than in concentrated saline solutions and, consequently, the vapor pressures of these solutes are much higher than those encountered over water at the same concentration. Therefore, the driving force for the vapor transfer of these solutes from the juice to the stripping solution is lower when compared with the driving force of the thermal evaporation. Finally, the diffusive permeabilities of these solutes through the membrane are lower than the water due to their higher molecular weights [34]. As a result, the loss of volatile compounds in OD can be markedly limited in comparison with the traditional evaporation and the organoleptic characteristics of the original juice are very well preserved. 8.3.2 OD Membranes and Modules

Membranes used in OD are typically hydrophobic in nature and realized with apolar polymers with low surface free energy such as PE, PTFE, PP and PVDF. However, OD membranes can also be realized by grafting on the surface of hydrophilic ceramic membranes molecules containing hydrophobic ﬂuorocarbon chains like ﬂuoroalkylsilanes or by coating the surface of alumina membranes with a thin lipid ﬁlm [37, 38]. The risk of wetting of the hydrophobic membrane, with a consequent reduction of the evaporation ﬂux and separation performance, is the main drawback of the OD process. Consequently, the membrane surface should be sufﬁciently hydrophobic in

8.3 Osmotic Distillation

order to prevent penetration of both feed and strip solutions into the pores by capillary forces (the contact angle between the liquid and solid phase should be greater than 90 ). Furthermore, the surface tension of the liquids should be high enough so that the capillary penetration pressure into the membrane pores is higher than the maximum operating transmembrane pressure difference in order to avoid a mixing of feed and strip solutions. For most concentrated salt solutions the surface tension is higher than that of pure water imploring the intrusion of such solutions into membrane pores at pressures normally used in the OD process. However, some fruit juices, especially citrus juices, contain peel oils and other highly lipophilic components that reduce their surface tension and promote wetting of hydrophobic surfaces. In this case, membrane materials with contact angles much higher than the water are needed. Furthermore, surface-active agents contained in cleaning solutions can also promote membrane wet-out and liquid penetration in OD membranes. Hydrophilic polymers, including cross-linked gelatin, agar, cross-linked polyacrylamide, esters of cellulose, cross-linked polyvinylalcohol, can be used to produce laminate membranes preventing liquid intrusion without impeding vapor transport [39]. The hydrogel-ﬁlm-side of the laminate should be in contact with the feed liquid or solution to be concentrated. Cellophane membranes and sodium alginate hydrogel coating on PTFE membranes has been also proposed to prevent membrane wet-out [40, 41]. Hydrophobic membranes resistant to oily feeds (e.g., limonene solution) have been developed by Mansouri and Fane [39] by coating the feed side of commercial ﬂat-sheet membranes, including Celgard 2500 (PP/PE, Hoechst Celanese), Durapore GVPS (PVDF, Millipore) and the UPVP (ultrahigh molecular weight PE, Millipore), with a thin layer of polyvinylalcohol (PVA). The laminate membranes were stable in oil emulsions for concentrations up to 1 wt.% for periods up to 24 h; on the contrary, membranes without the hydrogel coating were wetted out very rapidly by the oily feed. Membranes for OD applications should be highly porous (60–80%) and as thin as possible (0.1–1 mm) since the ﬂux is directly proportional to the porosity and inversely proportional to the membrane thickness. Basically, the overall thickness for OD membranes can vary from 80 to 250 mm, depending on the absence or presence of support. The thermal conductance of an OD membrane should be sufﬁciently high so that the energy of vaporization can be supplied by conduction across the membrane at a low temperature gradient. Consequently, the temperature difference between the two sides of the membrane is quite small (generally not exceeding 2 C), making the process isothermal. Barbe et al. [20] found higher organic volatile retentions per unit water removal in membranes characterized by large pore size at the surface when compared with membranes with small surface opening. These membranes offered a greater intrusion of the feed and the stripping solution, resulting in an increase in the resistance of the boundary layer to the diffusion of volatile components. Mengual et al. [31] found also that membranes with a small pore size (0.05–0.5 mm) did not show a signiﬁcant change in the transmembrane ﬂux. A similar behavior was also observed by Brodard et al. [42] in ceramic membranes with pore sizes in the range 0.2–0.8 mm.

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Stirred membrane cell, plate and frame, tubular, spiral-wound and hollow ﬁber are typical membrane modules used for OD operations. Flat-sheet membranes are preferred for pilot-scale studies for their versatility when compared with tubular and hollow-ﬁber membrane modules. However, many studies on the pilot scale have been performed by also using tubular, hollow ﬁber and spiralwound membranes. Plate-and-frame modules with a net-shaped spacer on the extract side and a smooth juice-side path have also been developed for the concentration of unclariﬁed juice with a high pulp content [43]. Helically wound hollow-ﬁber modules offer an improvement in the hydrodynamic conditions on the shell-side if compared with axial ﬂow modules; consequently, higher concentration of solutes and higher evaporation ﬂuxes can be obtained when viscous feeds are processed [44]. The most well-known module designed for OD is the Liqui-CelÒ Extra-Flow membrane contactor (Membrana-Charlotte, North Carolina, USA). It is constituted by microporous polypropylene hollow-ﬁber membranes approximately 300 mm in external diameter with a mean pore diameter of about 30 nm and a porosity of about 40%. The ﬁbers are potted into a polyethylene tubesheet and the shell casing is polypropylene, PVDF or 316L stainless steel [27, 34, 45]. The smallest modules are 2.5 inches in diameter with a membrane surface area of 1.4 m2, while the largest are 10 inches in diameter and offer a contact area of 130 m2. Commercial asymmetric OD membranes characterized by a thin PTFE layer supported by PP net have been manufactured by Pall-Gelman (East Hills, NY, USA). The top layer offers a resistance to the gas transfer, while the membrane support offers an additional resistance to water transfer in the liquid form. The mass-transfer resistance in the vapor phase is about 40–70% of the total resistance. The resistance of diluted brine entrapped in the PP support can cover up to 30% of the total resistance and the diluted brine boundary layer up to 60%, indicating the sensitivity of the OD system to concentration polarization phenomena [33]. 8.3.3 Effect of Operating Conditions on the OD Flux

The water-vapor ﬂux in OD is affected by different operating conditions. First, the OD ﬂux is signiﬁcantly affected by the solute content of the stripping solution. In particular, an increase of the transmembrane ﬂux by increasing the concentration of the stripping solution was observed both in real systems [46, 47] and in model systems in which water was used as feed [48–50]. These results may be explained assuming the strong dependence of the water activity of the stripping solution on salt content. The OD ﬂux is also differently affected by the type of stripping solution. Nagaraj et al. [46] studied the concentration of pineapple juice by OD by using calcium chloride and sodium chloride as stripping solutions. Calcium chloride produced higher ﬂuxes due to its higher osmotic activity, which resulted in a higher vapor pressure gradient across the membrane. Celere and Gostoli [51] compared the evaporation ﬂuxes in OD by using aqueous solutions of propylene glycol, glycerol

8.3 Osmotic Distillation

and glycerol–salt mixtures as an alternative to calcium chloride in order to overcome the problem of corrosion and scaling associated with the use of brines. Propylene glycol and glycerol solutions (70–75 wt.%) resulted less effective than highly concentrated CaCl2 and exhibited a similar extractive power. Ternary mixtures water–glycerol–NaCl were characterized by lower viscosities in comparison with the glycerol alone and offered similar ﬂuxes. The OD ﬂux is also affected by the feed concentration. Ravindra Babu et al. [47] reported a decreasing of the evaporation ﬂux in the concentration of sweet-lime juice and phycocianin solution by OD when the feed concentration is raised. A similar behavior was observed by Sheng et al. [52] during the concentration of apple, orange and grape juice through a PTFE membrane and by Courel et al. [49] when sugar solutions of increasing sucrose content were dehydrated by using stripping solutions of 45.5 w/w% initial CaCl2 content. This phenomenon can be attributed to the exponential increase of the viscosity and a decrease of the diffusion coefﬁcient when the solute content is raised. The increasing viscosity results in an increase of the concentration polarization effect, which reduces the driving force and, consequently, the evaporation ﬂux [53]. The transmembrane ﬂux increases by increasing the ﬂow rate of the osmotic agent due to the reduction of the concentration polarization layer along the condensation side of the membrane. This phenomenon was observed in the OD of pure water [33], sucrose solutions [49] and sweet-lime juice [46, 47]. Ravindra Babu et al. [54] observed also an increase in transmembrane ﬂux when the ﬂow rate of pineapple juice was increased from 25 ml/min to 100 ml/min; the increase in transmembrane ﬂux can be explained assuming a reduction in concentration polarization effect on the feed side. The increase in ﬂux, however, was more prominent (about 20%) by increasing the osmotic agent velocity. This phenomenon can be attributed to a lower concentration polarization on the feed side if compared to that on the brine side. Finally, the OD ﬂux is strongly affected by the feed temperature. Courel et al. [49] reported an increase of the evaporation ﬂuxes of 120% in the range 20–35 C for a 35 w/w% sucrose solution. Similarly, Bui and Nguyen [55] reported a 200% increase in the evaporation ﬂux for a feed temperature increasing of 20 C in the concentration of 40 and 50 w/w% aqueous glucose solutions by means of PVDF hollow ﬁbers. The increase in the OD ﬂux when the temperature is raised can be explained assuming an exponential type relation between the vapor pressure difference across the membrane and the temperature according to Clapeyrons law. Moreover, an increasing temperature determines a decrease in the feed and brine viscosities and an increasing of the solute diffusion coefﬁcient. In Figure 8.2 the general trend showing the effect of process parameters (concentration and ﬂow rate of feed and osmotic agent, operating temperature) on OD evaporation ﬂuxes is reported. Experimental curves are referred to the concentration of clariﬁed grape must by using a hollow-ﬁber polypropylene OD membrane module and calcium chloride dehydrate as stripping solution [56]. Vaillant et al. [57] evaluated the potential of OD for concentrating clariﬁed passion fruit juice on an industrial scale at 30 C, up to a total soluble solids (TSS) higher

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0.8

0.72

0.75

0.8 J Δa

0.6

0.64

0.6

0.56

0.4

0.48

0.2

0.40 70

0.0

0.50 0.25

0

10

20

30

40

50

60

Δaw

Δaw J (kg/m2h)

J (kg/m2h)

J Δa

0.00

1.0

0.80

1.00

0.4 0.2

0

10

20

30

50

60

0.0 70

[CaCl2 x2H2O] (%w/w)

TSS feed (°Brix) (a)

(b) 0.65

0.8

0.60

J (kg/m2h)

1.0

0.6 0.4

0.55 0.50 0.45

0.2

0.40 0.0

0.0 1

2

3

4

5

6

7

8

feed velocity (10-3 m/s)

0.5

1.0

1.5

2.0

2.5

Stripping solution velocity (10-2 m/s)

(c)

(d)

1.50

J (kg/m2h)

J (kg/m2h)

40

1.25

1.00

0.75

0.50 20

25

30

35

40

45

T (°C) (e) Figure 8.2 Concentration of clarified grape must by OD. Evaporation flux (J) as a function of: (a) grape must concentration; (b) stripping solution concentration; (c) flow rate of grape

must at a TSS content of 19.2 Brix; (d) flow rate of stripping solution (calcium chloride dehydrate); (e) operating temperature. Reprinted from [56] with permission of Chiriotti Ed.

than 60 Brix, by using a pilot plant containing a 10 m2 hollow-ﬁber module. Average evaporation ﬂuxes of 0.65 kg m2 h1 and of 0.50 kg m2 h1 were obtained at 40 and 60 Brix, respectively. These values were 10 times lower than those obtained in RO. The ﬂux decay during OD was attributable more to the dilution of

8.3 Osmotic Distillation

the stripping solution at a low TSS of the juice, while it mainly depended on the juice viscosity when juice concentration reached a value higher than 40 Brix. Cassano et al. [58–60] conﬁrmed these observations in the concentration of clariﬁed kiwifruit, orange and cactus-pear juice, by using a Liqui-Cel membrane module containing PP hollow ﬁbers. 8.3.4 OD Applications

Most OD applications in fruit-juice processing are related to the concentration of fruit juices up to concentrations of total soluble solids higher than 60 Brix, values signiﬁcantly higher than those achievable by RO. OD has the ability to remove selectively water from low molecular weight compounds, both volatiles and nonvolatiles, producing concentrated juices having superior quality. However, it is characterized by low ﬂuxes and it is inherently more costly when compared to thermal evaporation and RO [34]. In order to overcome these drawbacks and to improve the economics, many researchers have developed integrated membrane operations involving the clariﬁcation of the initial raw juice (eventually after a depectinization process), an optional preconcentration of the clariﬁed juice and a ﬁnal concentration of the clariﬁed or preconcentrated juice by OD. Table 8.1 summarizes OD applications concerning the concentration of fruit and vegetable juices including membrane type, stripping solution and type of treated juice. Few studies refer to the treatment of unclariﬁed juices. In particular, unclariﬁed noni juice was concentrated from 8 up to 32 Brix using a CaCl2 solution with an initial concentration of 6 mol kg1 as extraction brine. At isothermal conditions (30 C) transmembrane water-vapor ﬂuxes ranged between 0.09 and 0.413 kg m2 h1. Phenolic compounds were well preserved during the concentration step [61]. Ravindra Babu et al. [47] studied the inﬂuence of the osmotic agent concentration, ﬂow rate of feed and osmotic agent and membrane pore size on transmembrane ﬂux in the OD concentration of unclariﬁed sweet-lime juice. The juice was concentrated from 5 to 55 Brix at ambient temperature and atmospheric pressure. The masstransfer mechanism was in the transition region between Knudsen and molecular diffusion. The contribution of Knudsen diffusion was higher when the membrane pore size was 0.05 mm, while for a pore size of 0.20 mm molecular diffusion was the prevailing mechanism. The retention of volatile organic ﬂavor/fragrance components in the concentration of both unclariﬁed Gordo grape juice and Valencia orange juice was evaluated by using two ﬂat-sheet OD membranes (Celgard 2500 and Goretex L31189) of different materials (polypropylene and PTFE) and pore diameters at the surface (0.27 and 1.08 mm) by using 45% CaCl2 as a stripper [20]. As in the case of model solutions the degree of organic volatiles retention was the greatest for the membrane with the largest pore diameter at the surface (Goretex L31189). The clariﬁcation step performed by MF or UF allows removal of suspended solids and colloids minimizing possible fouling of either RO or OD units. UF of single-

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Table 8.1 OD applications in fruit-juice concentration.

Fruit juice

Membrane type

Stripping solution

Reference

Noni

Liquicel, hollow ﬁber, polypropylene Accurel Enka, ﬂat-sheet, polypropylene

Calcium chloride, 2, 4 and 6 M Calcium chloride dehydrate, 2–14 M Sodium chloride, 2–6 M Calcium chloride dehydrate, 2, 4, 6, 8 and 10 m Calcium chloride dehydrate, 2, 4, 6, 8, 10 m Sodium chloride, 2, 3, 4, 5 and 6 m Calcium chloride dehydrate, 10–60% w/w Calcium chloride dehydrate, 6 M

[61]

Calcium chloride, 4.6 M

[63]

Calcium chloride, 4.6 M

[64]

Calcium chloride dehydrate, 60% w/w Calcium chloride, 45% w/w

[59, 65, 66]

Pineapple and sweet lime (clariﬁed by pectinase) Pineapple (clariﬁed by bentonite) Sweet-lime

Accurel Enka, ﬂat-sheet, polypropylene Accurel Enka, ﬂat-sheet, polypropylene

Grape (clariﬁed by UF)

Liquicel, hollow ﬁber, polypropylene Microdyn, tubular, polypropylene

Blackcurrant, redcurrant, sour cherry, raspberry (clariﬁed by UF) Orange and passionfruit (clariﬁed by MF) Pineapple (clariﬁed by MF) Kiwi (clariﬁed by UF) Orange and grape

Raspberry, sour cherry, redcurrant, blackcurrant Camu-camu (clariﬁed by MF) Chokeberry, redcurrant, cherry (clariﬁed by UF) Cactus pear (clariﬁed by UF) Citrus and carrot (clariﬁed by UF and preconcentrated by RO) Red orange (clariﬁed by UF and preconcentrated by RO) Apple Grape (clariﬁed by UF) Anthocyanin extract from red radishes (clariﬁed by UF and preconcentrated by RO)

Hollow ﬁber, polypropylene Hollow ﬁber, polypropylene Liquicel, hollow ﬁber, polypropylene Celgard 2500, ﬂat-sheet, polypropylene Goretex L31189, ﬂat-sheet, PTFE Microdyn, tubular, polypropylene Pall-Gelman TF200, ﬂatsheet, PTFE Microdyn, tubular, polypropylene Liquicel, hollow ﬁber, polypropylene Liquicel, hollow ﬁber, polypropylene

[46]

[54] [47]

[56] [62]

[20]

Calcium chloride, 6 M

[67]

Calcium chloride, 4.0–5.2 M Calcium chloride dehydrate, 6 M Calcium chloride dehydrate, 60% w/w Calcium chloride dehydrate, 4.1–4.5 M

[68] [69] [60] [58]

Liquicel, hollow ﬁber, polypropylene

Calcium chloride dehydrate, 60% w/w

[70]

Microdyn, tubular, polypropylene Liquicel, hollow ﬁber, polypropylene Flat-sheet, polypropylene

Calcium chloride, 3.5 and 6M Calcium chloride, 40% w/w Calcium chloride dehydrate, potassium phosphate

[62, 67] [71] [72]

8.3 Osmotic Distillation Table 8.1 (Continued)

Fruit juice

Membrane type

Stripping solution

Reference

Passionfruit (clariﬁed by MF) Orange juice (diluted from commercial concentrate) Pineapple (unclariﬁed and clariﬁed by MF) Orange (clariﬁed by MF)

Hollow ﬁber, polypropylene Accurel, follow ﬁber, polypropylene

Calcium chloride, 45%w/w Calcium chloride dehydrate, 4.9 M

[57]

Pall-Gelman TF200, ﬂatsheet, PTFE Hollow ﬁber, polypropylene Hollow ﬁber, polypropylene Microdyn, hollow ﬁber, polypropylene

Calcium chloride, 5.5–6.0 M Calcium chloride, 5.5 M

[74]

Calcium chloride, 5.3–5.6 M Calcium chloride dehydrate, 65 Brix

[76]

Melon (clariﬁed by MF) Blackcurrant (clariﬁed by MF and preconcentrated by RO)

[73]

[75]

[77]

strength Gordo grape juice using membranes with nominal pore diameters of 0.1 mm or less resulted in appreciable osmotic distillation ﬂux increases over that observed for juice not subjected to the UF [71]. Similarly, evaporation ﬂuxes between 7 and 10 kg m2 h1 were obtained by Hongvaleerat et al. [74] in the concentration of clariﬁed pineapple juice. These values were higher than those obtained with the single-strength juice due to the complete removal of suspended solids in the clariﬁcation step. This phenomenon can be attributed to a reduction in the viscosity of the concentrated juice-membrane boundary layer where the solute concentration is highest and the effect of protein removal is expected to be more pronounced. UF pretreatment results also in a small increase in juice surface tension with a consequent reduction in the tendency for membrane wet-out to occur. Rodrigues et al. [68] evaluated the performance of the OD and RO processes in the concentration of camu-camu juice previously clariﬁed by MF. RO permitted to reach higher ﬂuxes (50 kg m2 h1) than OD, but a lower concentration of soluble solids (25 Brix). OD allowed to concentrate the juice up to 63 Brix with evaporation ﬂux values of 10 kg m2 h1. Several studies demonstrated the efﬁciency of the OD process in maintaining the nutritional, sensorial and organoleptic characteristics of the original juice. An integrated MF/OD membrane process was implemented on a semi-industrial scale by Cisse et al. [75] to produce concentrated orange juices. The clariﬁed juice was concentrated through a two-stage OD process producing concentrated juices at 45 and 62 Brix, respectively. Most aroma compounds and ascorbic acid were recovered in the clariﬁed juice, while apolar compounds, such as terpenic hydrocarbons and carotenoids, were retained by the MF membrane. Signiﬁcant losses of Vitamin C (from 6 to 15%) were mainly observed at the beginning of the concentration. This phenomenon was attributed to the Vitamin C oxidation by the residual oxygen

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entrapped within the pores of the membrane. As the residual oxygen contained in the circuit was consumed, vitamin C losses decreased during processing. The low temperatures employed (lower than 28 C) in the concentration steps preserved the color of the juice (L-values, hue angle and color purity of both concentrates were similar to those of the clariﬁed juice). Also, the sugar and acid contents were not modiﬁed during the juice concentration. Losses of aroma compounds were higher in the ﬁrst OD stage (about 31%) than in the ﬁnal one (about 22%). However, sensorial analyses revealed no signiﬁcant differences between the clariﬁed juice and the ﬁnal concentrate juice at 62 Brix. Further, the quality of the pulpy juice obtained by mixing the MF retentate, previously pasteurized, with the OD concentrate was high and much closer to that of the initial single-strength juice than to the commercial thermal concentrate. An excellent preservation (>97%) of the total antioxidant activity (TAA) of red fruit juices (chokeberry, redcurrant and cherry juices) was also observed in an UF-OD sequence investigated by Koroknai et al. [69]. Integrated membrane processes were also studied and proposed by Cassano et al. [58–60, 65] to produce high quality concentrated fruit juices (such as kiwifruit, orange, lemon and cactus pear) with ﬁnal TSS concentration of 63–65 Brix. Fresh juices were previously depectinized, clariﬁed by UF and then optionally preconcentrated by RO before the ﬁnal OD concentration step. OD was performed by using a Liqui-Cel Extra-Flow 2.5 8 inch membrane contactor equipped with polypropylene hollow ﬁbers and a calcium chloride dehydrate solution at 60%w/w as the stripping solution. During the concentration process of red orange juice, a slight decrease of the total antioxidant activity (TAA) was observed (about 15%) due to the partial degradation of ascorbic acid (about 15%) and anthocyanins (about 20%). Nevertheless, this degradation was lower than that observed in the thermally concentrated juice where anthocyanins and hydroxycinnamates (particularly ferulic and p-coumaric acid) underwent a reduction of 36% and 55%, respectively; for the ascorbic acid and ﬂavonones removals were in the order of 30 and 23%, respectively [70]. Table 8.2 shows the physical characterization of the Chilean kiwifruit juice (Hayward variety) and of clariﬁed and concentrated fractions obtained in an integrated UF/OD process. The UF process determines a complete removal of suspended solids and turbidity in the depectinized raw juice. Most sugars are recovered in the clariﬁed juice; the low reduction of TSS content in the clariﬁed juice can be attributed to the removal of suspended solids that, together with the soluble pectin, can interfere with the measurements of the refractive index. In Table 8.3 measurements of total antioxidant activity (TAA) and ascorbic acid in clariﬁed and concentrated fractions are reported. In particular, the clariﬁed juice showed a reduction of vitamin C and TAA of 16 and 8%, respectively, in comparison with the fresh juice. During the OD process the vitamin C content was constant independent of the achieved TSS level. TAA of samples concentrated at 20 and 30 Brix was similar to that of the clariﬁed juice while a little reduction was observed at higher concentrations. On the contrary, the juice concentrated by thermal evaporation at 65 Brix showed an 87% reduction of the ascorbic acid when compared to the

8.3 Osmotic Distillation Table 8.2 Physical characterization of Chilean kiwifruit juice clarified and concentrated by integrated UF/OD process.

Parameter Total soluble solids ( Brix) Turbidity (NTU) Viscosity at 25 C (MPa s) pH Suspended solids (% w/w)

Feed

Permeate UF

Retentate OD

12.6 299.5 1.455 3.58 17.0

12.1 0 1.427 3.60 0

61.4 — 44.5 3.40 —

clariﬁed juice. The TAA was reduced by about 50% independent of the TSS content achieved [66]. A pervaporation (PV) step was also investigated to recover aroma compounds from kiwifruit juice and introduced in an integrated UF-OD process in order to evaluate the best conﬁguration giving the minimal loss of aroma compounds. For the majority of the aroma compounds detected, the enrichment factor in the permeate of the fresh juice was higher than the clariﬁed and concentrated juice suggesting the use of PV for the removal and enrichment of aroma compounds directly from the fresh juice, before any concentration process [65]. Shaw et al. [63] evaluated the retention of ﬂavors in concentrated orange and passionfruit juices (previously clariﬁed by MF) obtained by using a pilot-scale osmotic evaporator containing 10.3 m2 of PP hollow ﬁbers. Both juices were concentrated threefold to 33.5 and 43.5 Brix, respectively. Quantitative headspace gas chromatographic analyses showed a loss of volatile compounds of about 32% and 39% in orange and passionfruit juice, respectively. A multistep membrane process on laboratory and large scale was implemented by Kozak et al. [77] for the production of blackcurrant concentrated juices. The raw juice (15–18 Brix as TSS content) was clariﬁed by MF and then preconcentrated by RO. The ﬁnal concentration step was performed by using OD in which the TSS Table 8.3 Analyses of TAA and ascorbic acid in Chilean kiwifruit juice clarified and concentrated by integrated UF/OD process.

Sample

TSS ( Brix)

Ascorbic acid (g l1)

TAA (mM Trolox)

UF-F UF-P UF-R OD-R1 OD-R2 OD-R3 OD-R4 OD-R5

12.6 11.2 12.4 20.0 35.0 44.0 53.8 61.2

0.90 0.75 0.72 0.81 0.84 0.81 0.81 0.82

17.6 16.2 16.9 16.2 16.2 15.5 15.4 15.1

Legend: F: feed; P: permeate; R: retentate.

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content of the juice was increased up to 63–72 Brix. Results of sensory proﬁle analyses revealed that the color intensity, the transparent ability and the acidic ﬂavor intensity of the concentrated juice were similar to those of the raw juice. The anthocyanin content of the concentrated juice was more than three times higher than that of the raw juice. The effect of an integrated MF/OD process on the physicochemical, nutritional and microbiological qualities of melon juice was investigated by Vaillant et al. [76]. The raw juice was macerated with an enzymatic solution containing hemicellulase and cellulose activities before the MF step performed by using a ceramic multichannel membrane. Average permeation ﬂuxes of about 80 l m2 h1 were obtained with the continuous extraction of retentate at a volumetric reduction ratio of 3. OD was performed by using a module containing polypropylene hollow-ﬁber membranes circulating the cold clariﬁed juice (6 C) in the lumen of the ﬁbers, and calcium chloride brine 5.3–5.6 M in the shell side. The clariﬁed juice was concentrated from 7 to 55 Brix of TSS. Evaporation ﬂuxes decreased from 0.7 to 0.57 kg m2 h1 when juice TSS reached the ﬁnal value. Insoluble solids of the raw juice were removed in the MF step; the MF membrane also rejected b-carotene probably due to its association with membrane and wall structures of the cell fragments. Thus, the authors proposed to use the pulpy juice (retentate MF) as raw material to extract b-carotene or directly in functional drinks. MF also ensured the microbiological stability of the juice in a single operation. In the OD concentrate the acidity, the color and the sugar content of the clariﬁed were well preserved. No signiﬁcant loss of vitamin C was observed in comparison with the clariﬁed juice, while a loss in polyphenol compounds of about 30% was attributed to the presence of polyphenol oxidases in the clariﬁed juice still acting during juice concentration. Most OD studies in fruit-juice concentration refer to applications on laboratory scale. A successful application on pilot scale was conducted in pilot-plant facilities located in Mildura and Melbourne, Australia. The Melbourne facility, designed by Zenon Environmental (Burlington, Ont.), was a hybrid plant consisting of UFand RO pretreatment stages, an OD section containing two 19.2-m2 Liqui-Cel membrane modules and a single-stage brine evaporator. Fresh fruit juices were concentrated up to 65–70 Brix at an average throughput of 50 l h1. The system was used to develop operational parameters and economic data for the design of full-scale plants and for the production of concentrated samples for their testing and evaluation [34]. The Mildura plant, designed by Vineland Concentrates and Celgard LLC, contained 22 Liqui-Cel membrane modules (4 28 inches) for a total interfacial area of 425 m2. It was used for the concentration of grape juices to make wine from reconstituted concentrate. The installation, having a feed rate of approximately 80–100 1 h1 was able to produce approximately 20–25 1 h1 of 68 Brix concentrate [78]. In Figure 8.3 a general scheme for the production of fruit-juice concentrate by integrated membrane operations is reported. Concentrated juices of good quality, not exceeding TSS concentrations of 30 Brix, can be obtained by RO. In this case, the loss of volatile and nonvolatile ﬂavors can be minimized, limiting the fraction of water removed. The further concentration of the RO retentate by OD permits a concentrated product, similar to that achieved by using OD alone, to be obtained

8.4 Membrane Distillation Fruits

Fruithandling andjuiceextraction water (fruitcleaning, etc.)

Washing Milling RO

Pressing

PRECONCENTRATION FRUITJUICE enzymes

evaporator

diluted brine

UF

muds

CLARIFICATION

condensate condenser

brine

OD

CONCENTRATION

Figure 8.3 Process scheme of integrated membrane operations for the production of concentrated fruit juices.

reducing signiﬁcantly the processing costs. The stripping solution used in the OD step can be reconcentrated by thermal evaporation, cooled and recycled to the OD system. It has been suggested that the RO preconcentration of a fruit juice with a TSS content of 18 Brix to about half its original volume could reduce the amount of water needed to concentrate the juice by OD up to 68 Brix of about 56% maintaining unchanged its quality. Furthermore, a reduction of the evaporator capacity requirement and of the OD membrane area can be achieved, leading to a decrease in capital and operating costs [34]. The management of the diluted brine strep is one of the drawbacks associated with the commercial application of the OD in fruit-juice processing. Although the regeneration of exhausted brines can be realized by thermal evaporation, this operation is expensive due to corrosion and scaling phenomena. Solar ponding, RO and pervaporation have been proposed as alternatives to reconcentrate the diluted brine solutions [78]. Electrodialysis has been also suggested for the regeneration of NaCl brines [79]. 8.4 Membrane Distillation 8.4.1 Process Fundamentals

Membrane distillation (MD) is a thermally driven process in which two aqueous solutions, at different temperatures, are separated by a microporous hydrophobic membrane to support a vapor/liquid interface. The hydrophobic nature of the membrane prevents liquid solutions from entering its pores due to the surface-tension forces. In these conditions, a water-vapor transfer from the warm

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side to the cold one occurs. The driving force of the process, as in the OD operation, is a vapor-pressure difference between the two solutions separated by the membrane; however, unlike the OD process it is generated by a temperature gradient rather than a concentration gradient. The process takes place at a temperature that may be much lower than the boiling point of the solutions (feed temperatures in MD typically range from 60 to 90 C, although temperatures as low as 30 C have been used). Operating pressures are generally of few hundred kPa, relatively low if compared to pressure-driven processes such as RO. Consequently, equipment costs and mechanical demands on the membrane are greatly reduced. These features make MD ideal for the treatment of food and pharmaceutical solutions. Unlike pressure-driven processes, MD membranes operate as a support for a vapor/liquid interface and not as sieving devices. Therefore, they can be realized with chemically resistant polymers such as PTFE, PP and PVDF. Furthermore, fouling phenomena are greatly reduced since the pores are relatively large if compared to the pores or to the diffusional pathway in RO or UF and are not easily clogged. Finally, MD membranes operate on the principles of vapor–liquid equilibrium so that nonvolatile compounds (such as ions, colloids, macromolecules, cells, etc.) are totally rejected [80]. Basically, a vapor-pressure difference across a MD membrane can be realized through four different methods. In the direct contact membrane distillation (DCMD) the permeate side of the membrane consists of a condensing ﬂuid in direct contact with the membrane (cold distillate) separated by the hot feed. In this case, the vaporpressure gradient that results from the transmembrane temperature difference is the driving force of the mass transport across the membrane. DCMD is the most suitable technique for applications in which the volatile component is water. The water transport through the membrane can be summarized in three steps: (1) evaporation of water at the dilute vapor/liquid interface; (2) diffusion or convective vapor transport through the membrane pores; (3) condensation of water vapor. In the air-gap membrane distillation (AGMD) the permeate side of the membrane is a condensing surface separated from the membrane by an air gap. Volatile molecules cross both the membrane pores and the air gap and ﬁnally condense over a cold surface inside the membrane module. In the sweeping gas membrane distillation (SGMD) a cold inert gas sweeps the permeate side of the membrane carrying the volatile molecules. In this case, condensation occurs outside the membrane module [81, 82]. Finally, in the vacuum membrane distillation (VMD) vacuum is applied on the permeate side of the MD membrane by means of a vacuum pump and, similarly to SGMD, condensation takes place outside the membrane module [83, 84]. 8.4.2 MD Membranes and Modules

Microporous membranes for MD are typically realized in ﬂat-sheet or tubular form with hydrophobic polymers such as polytetraﬂuoroetilene (PTFE), polypropylene

8.4 Membrane Distillation

(PP), polytehylene (PE) and polyvinylidene diﬂuoride (PVDF). They can be prepared by phase inversion, stretching of dense ﬁlm and thermally induced phase separation. Hydrophilic membranes can be also used for MD applications after treatments aiming to make their surfaces hydrophobic [85–87]. The transmembrane ﬂux through a MD membrane is related to the membrane pore size and other characteristic parameters by the following equation: N/

rae dm t

ð8:4Þ

where N is the molar ﬂux, r the mean pore size of the membrane pores, a a factor whose value is 1 for Knudsen diffusion and 2 for viscous ﬂuxes, respectively, dm the membrane thickness, e the membrane porosity and t the membrane tortuosity [80, 88]. According to Equation 8.4 the thinner the membrane and the greater the porosity of the membrane, the greater the ﬂux rate. On the contrary to achieve better heat efﬁciency the membrane should be as thick as possible in order to limit heat loss by conduction through the membrane matrix [89]. Typical pore sizes of MD membranes are in the range 0.2–1.0 mm. Membrane pores should be large enough to facilitate the required ﬂux; however, in order to prevent membrane pore wettability, the pore size should be as small as possible. Schneider et al. [90] recommended a maximum pore radius of 0.5–0.6 mm in order to avoid membrane wetting due to ﬂuctuations in process pressure and temperature. MD membranes should also be characterized by low thermal conductivity to prevent heat loss through the membrane matrix, good thermal stability and chemical resistance towards different feed solutions. In general, the porosity of MD membranes ranges between 30 and 85% of the volume and the overall thickness from 80 to 250 mm, depending on the absence or presence of support. Typical MD membranes show a pore-size distribution rather than a uniform pore size. Consequently, different mechanisms can occur simultaneously depending on the pore size and on the MD operating conditions. The design of MD modules should provide high feed and permeate ﬂow rates with high turbulence and low pressure drop along the membrane module. Good heat recovery function, thermal stability and high packing density should be also guaranteed. Flat-sheet membranes assembled in plate and frame or spiral-wound modules and capillary membranes in tubular modules have been used in MD operations. Most membrane modules for MD applications on the laboratory scale are assembled with ﬂat-sheet membranes since they offer a higher versatility than tubular or hollow ﬁber conﬁguration. Flat-sheet membranes can be easily removed from their modules for cleaning and examination; furthermore, the same membrane module can be used to evaluate the performance of different MD membranes. On the other hand, tubular membranes offer a much higher membrane surface area to module volume ratio and they are preferred for industrial applications. Additionally, they do not require a support, reducing the boundary-layer resistance in comparison with ﬂat-sheet modules.

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One of the ﬁrst commercially available membrane module for MD operations was realized by Enka AG (Akzo) in a shell-and-tube conﬁguration. On the industrial scale ﬂat-sheet modules have been produced by Gore and Associates (in a spiral-wound conﬁguration) and by the Swedish Development Co. (in a plate and frame conﬁguration). At the moment the availability of industrial membrane modules for MD applications is one of the main limitations for MD implementation. In most cases, the performance of the MD process is affected by the heat and mass transfer in the boundary layers. Furthermore, mass-transfer rates within boundary layers should be able to prevent excessive concentration polarization phenomena with consequent risks of wetting and scaling. 8.4.3 Effect of Operating Parameters on MD Fluxes

Permeate ﬂuxes in MD are affected by different operating conditions such as: feed concentration, operating temperature, feed circulation velocity, temperature difference, permeate inlet temperature, permeate ﬂow velocity and vapor-pressure difference. Permeate ﬂuxes decrease by increasing the feed concentration. This phenomenon can be attributed to the reduction of the driving force due to the decrease of the vapor pressure of the feed solution and to an exponential increase in the viscosity of the feed solution. The increase in the feed concentration also contributes to the formation of a boundary layer on the feed membrane surface (concentration polarization). However, this contribution is very small if compared to that of the temperature polarization [91]. At high concentration, ratio ﬂuxes observed in MD are higher than those observed in pressure-driven membrane processes [92, 93]. In all MD conﬁgurations the MD ﬂux increases exponentially with the increase of the feed temperature. This is due to the increase of the vapor pressure of the feed solution with temperature with a consequent increase of the transmembrane vapor pressure. However, when the treated aqueous solutions contain volatile compounds, such as fruit juices, the exponential increase of the permeate ﬂux with the feed temperature can be impeded by the drop in selectivity. This phenomenon can be attributed to the increased effect of both temperature and concentration polarization when the feed temperature is raised. The increase in feed circulation velocity and feed stirring rate determines an increase of the MD permeate ﬂux. This is due to the increase of the heat-transfer coefﬁcient in the feed side of the membrane module and to the reduction of the temperature and concentration polarization effects. The shear forces generated at high ﬂow rate cause a lower accumulation of particulates on the active membrane surface, thus reducing membrane fouling. Lower crossﬂow velocities cause a lower Reynolds number, thus preventing the heat transfer from the bulk of the solution to and from the membrane surface with an increase of the temperature polarization [94]. Consequently, higher performances in terms of productivity can be obtained by operating under a turbulent ﬂow regime. In some MD applications, an increase in

8.4 Membrane Distillation

ﬂux with the feed circulation velocity up to an asymptotic value at higher ﬂow rates is observed [95]; other studies report a linear increase of the MD ﬂuxes with the feed circulation velocity [96]. The temperature difference is another parameter affecting MD ﬂuxes. When the mean temperature is kept constant, a linear increase of the permeate ﬂux is observed by increasing the temperature difference [89]. On the other hand, the permeate ﬂux increases exponentially with the mean temperature when the temperature difference is maintained ﬁxed [31]. In DCMD applications, the general effect of increasing the permeate temperature is to reduce the permeate ﬂux: this can be attributed to the decrease of the transmembrane vapor pressure as long as the feed temperature is maintained constant [97]. The increase of the permeate velocity increases the heat transfer in the permeate side with a consequent reduction of temperature and concentration polarization effects. When the heat-transfer coefﬁcient in the permeate side is increased, the temperature at the membrane surface approaches the temperature in the bulk permeate side. This results in an increase in the driving force of the process and, consequently, of the permeate ﬂuxes. Finally, in all MD conﬁgurations a linear increase of the permeate ﬂux is obtained by increasing the transmembrane vapor-pressure difference between the two sides of the membrane. 8.4.4 MD Applications

Table 8.4 shows some selected MD applications in fruit-juice processing: they refer to the concentration of clariﬁed (or clariﬁed and preconcentrated) juices and to the recovery of aroma compounds by using DCMD and VMD conﬁgurations. The potentiality of MD in orange juice concentration by integrated membrane systems was analyzed by Drioli et al. [92] and Calabrò et al. [101] considering the effect of the viscosity and the necessity of juice pretreating. In these studies commercial plate PVDF membranes made by Millipore Corp., with a nominal pore radius of 0.11 mm and a porosity of 75%, were used for the concentration of single-strength orange juice with a TSS content of 10.8 Brix. The permeate ﬂux decreased from 5.3 l m2 h1 to about 2.5 l m2 h1 when the juice was concentrated up to 31 Brix at a transmembrane temperature gradient of 20 C (feed temperature ¼ 40 C; cooling water ¼ 20 C). Results on juice composition showed a very good retention of soluble solids, sugars and organic acids with rejection of sugars and organic acids equal to 100%. The observed reduction of vitamin C of about 42% was associated to high temperature and oxidation. The color and ﬂavor of the concentrated juice were satisfactory. The pretreatment of the juice by UF permitted removal of pulp and pectin and a clariﬁed juice with a lower viscosity compared with the single-strength juice to be obtained. An increase of the permeate ﬂux in MD was observed when the clariﬁed juice was concentrated from 10 to 40 Brix without ﬂux decay.

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Table 8.4 MD applications in fruit-juice processing.

Fruit juice

Membrane type

MD configuration

Reference

Apple (diluted from commercial concentrate) Apple (clariﬁed by UF) Blackcurrant (clariﬁed by UF) Blackcurrant (clariﬁed by UF and preconcentrated by RO) Orange juice (diluted from commercial concentrate)

Enka Microdyn, hollow ﬁber, polypropylene MFK3, ﬂat sheet, PVDF K150, ﬂat sheet, PTFE

DCMD

[97]

DCMD VMD

[98] [99]

Hollow ﬁber, polypropylene

DCMD

[100]

Millipore, ﬂat sheet, PVDF; Gelman, G0712, ﬂat sheet; Enka, hollow ﬁber, polypropylene Enka-Mycrodin, MD020 TP 2N, hollow ﬁber, polypropylene Akzo-Nobel, Accurel V8/ 2, tubular Enka Microdyn, hollow ﬁber, polypropylene

DCMD

[91, 101]

VMD

[102]

VMD

[103]

DCMD

[104]

Pear (model solution)

Must (model solution) Apple (diluted from commercial concentrate)

Legend: DCMD: direct contact membrane distillation; VMD: vacuum membrane distillation.

Kozak et al. [100] evaluated the applicability of MD to produce concentrated blackcurrant juice starting from the clariﬁcation of the raw juice by MF followed by a preconcentration step performed by RO. The preconcentrated juice, with a TSS content of 22 Brix, was concentrated up to 58.2 Brix by using a laboratory-size hollow-ﬁber polypropylene membrane module. The juice side and the water side temperature were maintained at 26 C and 11 C, respectively (DT ¼ 15 C). In these conditions a steady-state ﬂux of 0.45 kg m2 h1 was reached after the stabilization of the temperature. An increase of few degrees centigrade in the driving force (DT ¼ 19 C) increased the steady-state ﬂux to 0.8 kg m2 h1, markedly reducing the operation time required to concentrate the same amount of the juice. All the analyzed parameters (density, total acidity and anthocyanin content) were directly proportional to the increase of the TSS in the juice. Gunko et al. [98] applied the DCMD to the concentration of apple juice. The raw juice was depectinized and then submitted to an UF clariﬁcation step. The concentration of the clariﬁed juice was carried out by using PVDF microﬁltration membranes of MFFK3 type (NPO Polymersintes, Russian Federation) with a nominal pore size of 0.45 mm and a porosity of 80–85%. A TSS content of 50 Brix was obtained when the permeate ﬂux reached about 9 l m2 h1. Further concentration to 60–65 Brix resulted in reduced productivity (up to 3 l m2 h1). The permeate ﬂux increased on decreasing the cooling water temperature, maintaining constant the juice temperature in the hot cell.

8.4 Membrane Distillation

Highly concentrated apple juices up to 64 Brix were also produced by using polypropylene hollow-ﬁber DCMD modules (Enka MD-020-2N-CP) with tube and shell conﬁguration [97, 104]. Transmembrane ﬂuxes of about 1 kg m2 h1were obtained. Flux rates were dependent essentially upon temperature polarization phenomena located mainly on the feed side, rather than concentration polarization. Osmotic effects on the permeate side, obtained by using cold solutions of CaCl2, improved transmembrane ﬂuxes up to 20%. Model simulations describing the ﬂuid dynamics and the membrane behavior within the DCMS system were in good agreement with experimental results. In particular, a nonsymmetrical distribution gave results better than those obtained with a Gaussian distribution. An optimal membrane thickness value, ranging between 30 and 60 mm, was obtained from the model. The technical feasibility of must concentration through vacuum membrane distillation (VMD) was evaluated by Bandini and Sarti [103] with the object of increasing the alcoholic potential of musts, while preserving quality and quantity of the original aroma compounds. Experimental tests were performed on bench-scale plant by using model aqueous mixtures containing glucose and typical must aromas (1-hexanol, linalool, geraniol). A tubular membrane module containing Accurel V8/2 membranes (1.5 mm thickness, 5.5 mm internal diameter, 0.2 mm average pore size), manufactured by Akzo-Nobel and arranged in tube and shell conﬁguration was used. The aqueous mixture containing aromas was continuously recirculated in the tube side while the permeate vapors were removed from the shell side using a vacuum pump and condensed in cold traps refrigerated by liquid nitrogen. Experimental curves representing the transmembrane ﬂux against the concentration of the aqueous mixture at different values of feed temperatures and vacuumside pressures showed the typical trend encountered also in other VMD applications involving volatile organic compounds [105]. In particular, water ﬂux decreased by increasing the downstream pressure and the feed concentration, whereas an increasing of the liquid phase temperature determined an increase in the transmembrane ﬂux. On the basis of experimental results, a model to design a must concentration system was presented. The proposed system was an RO/VMD integrated process considering the relevant amount of water to be removed. The must is previously concentrated from 20 to 30 Brix in the RO unit and then submitted to a VMD step designed for the residual concentration up to 50 Brix at 60 C and 30 mbar. The proposed process allows production of juice concentrates containing at least 43% of the aroma originally contained in the fresh juice. Bagger-Jørgensen et al. [99] evaluated the potential of VMD in the recovery of aroma compounds from blackcurrant juice. Before the concentration step the raw juice was depectinized, clariﬁed with gelatin-silica sol, centrifuged and ﬁnally ultraﬁltered. VMD was performed at low temperatures (10–45 C) and at different feed ﬂow rates (100–500 l h1) by using a ﬂat PTFE (K150) membrane with a pore size of 0.1 mm manufactured by Osmonics. The highest concentration factors for the blackcurrant aroma compounds (from 21 to 31) were obtained at high feed ﬂow rate (400 l h1) and low temperatures (10 C). At 5 vol.% feed volume reduction the recovery of highly volatile compounds ranged

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between 68 and 83 vol.% and between 32 and 38 vol.% for the poorly volatile compounds. The recovery of aroma compounds by VMD from a mixture containing water/ ethanol/ethyl 2,4-decadienoate simulating the pear aroma compounds was also investigated by Diban et al. [102]. VMD experiments were performed by using a polypropylene hollow-ﬁber membrane module supplied by Enka-Mycrodin (unit MD 020 TP 2N) (pore diameter 0.2 mm, porosity 75%, inner diameter of hollow ﬁbers 5.5 mm). The highest values of enrichment factor (upto 15) for pear aroma compounds were obtained working at lower temperatures and higher downstream pressures. 8.4.5 Coupled Operation of Osmotic Distillation and Membrane Distillation

As previously described, both MD and OD processes are based on the use of similar membrane materials and module conﬁgurations and in both processes the driving force is a water-vapor pressure difference applied between the two sides of a hydrophobic membrane. The water transport through the membrane can be summarized in three steps including vaporization and condensation on the boundary layers and diffusion or convective vapor transport through the membrane pores. The simultaneous vaporization/condensation phenomenon at the liquid/membrane interface determines a temperature variation on both sides of the membrane: the stripping solution is heated by the latent heat of the condensation while the feed solution is cooled by the vaporization. This thermal effect, reducing the driving force of the water transport across the membrane, can be exploited to obtain a coupled process where the stripper and the feed solution are thermostated separately at different temperatures: the osmotic solution on the cold side and the solution to be concentrated on the warm side [62]. This coupled operation of MD and OD, referred to as membrane osmotic distillation (MOD), permits enhancement of the water ﬂux across the membrane [106]. The water transport through a porous hydrophobic membrane under a coupled transmembrane temperature/concentration difference was investigated by Godino et al. [107] in different experimental conditions (solute concentration, stirring rate, mean temperature and bulk temperature difference) by using pure water and sodium chloride solutions at different concentrations. A coupled MD/OD process for the concentration of fruit juices (apple, redcurrant and blackcurrant, sour cherry and raspberry) was studied by Koroknai et al. [67] by using a membrane contactor containing 34 polypropylene tubular membranes (Microdyn) with a nominal pore size of 0.2 mm, 70% porosity and thickness of 0.2 mm. The enzyme-treated raw fruit juice was pumped through the shell side, while the osmotic solution (CaCl2 6 M) was circulated in the lumen of ﬁbers in a countercurrent mode. Concentrated juices at 60 Brix were obtained in an operation time of 15–20 h maintaining a temperature difference on both sides of the membrane of 15 C. Enhanced water ﬂuxes were obtained since the total driving force of the process was higher than the sum of the driving forces of single processes. During the process the water ﬂux decay through the membrane as well as fouling phenomena

8.5 Supported-Liquid Membranes 2.5

J (kg/m2h)

2.0

1.5

1.0

0.5

0.0 -20

-15

-10

-5

0

5

10

15

20

ΔT (°C) Figure 8.4 Concentration of clarified grape must by OD. Effect of a transmembrane temperature gradient on the evaporation flux (J) along with a concentration gradient (stripping

solution ¼ calcium chloride dehydrate 60%, w/w). Reprinted from [56] with permission of Chiriotti Ed.

were not signiﬁcant. Furthermore, the organoleptic evaluation of the concentrated juices conﬁrmed their high quality when compared with commercial juices prepared from concentrates. Similarly, Belaﬁ-Bakó and Koroknai [108] found that the coupled process is more effective than MD or OD alone. The use of short membrane modules in cascade series, with heat exchangers placed between them, was suggested in order to minimize heat losses. Figure 8.4 shows the effect of the coupled transmembrane temperature/concentration gradient on the OD evaporation ﬂux obtained in the concentration of clariﬁed grape must by using a polypropylene hollow-ﬁber membrane module and calcium chloride dehydrate (60% w/w) as stripping solution. The temperature difference applied between the feed and the extracting solution varied from 15 C to þ 15 C, while the average temperature of the system was kept at 30 C. The increasing in OD ﬂux with a temperature gradient of þ 5 C, þ 10 C and þ 15 C, in comparison with the isothermal conditions, was of 53, 89 and 138%, respectively [56].

8.5 Supported-Liquid Membranes 8.5.1 Process Fundamentals

A supported-liquid membrane is a three-phase system in which a microporous support containing a carrier and a solvent is in contact with two aqueous solutions: the feed phase and the strip phase, respectively. The carrier binds the analyte of

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interest on the feed side, transfers it selectively through the micropores and releases it on the strip side. The microporous support is used to stabilize the carrier–solvent mixture that is held in the pores of the membrane due to capillary and surface forces [109]. If the membrane used is hydrophobic, an organic solution containing a carrier selective for one of the species in the feed solution is immobilized in the micropores. Hydrophilic membranes can also be used for SLM systems: in this case the carrier-containing solution is in aqueous phase. The selective transport is guaranteed if the afﬁnity between the compound of interest and the carrier is higher than that between the carrier and other compounds in the feed solution. The membrane pore size has to be large in comparison with the solute dimension in order to permit unhindered diffusion. Further, if the volumes of stagnant strip solutions are much smaller than the feed solution volumes, high preconcentration factors are achieved [110, 111]. The performance of the process is affected by the properties of the solution immobilized into micropores, such as viscosity and volatility, the carrier concentration and selectivity, the membrane characteristics and the ﬂuid dynamics of the aqueous phases. Basic aspects related to the stability of SLMs were reviewed by Kemperman et al. [112]. A rigorous model for mass transport through SLMs was proposed by Alhusseini and Ajbar [113]. 8.5.2 SLMs Applications

The main applications of SLMs include the separation of metals from hydrometallurgical wastewaters [114], fructose extraction from mixtures of sugars contained in fermentation broths [115], oleﬁn and parafﬁn hydrocarbon mixtures separation [116]. In the agro-food sector SLMs offer interesting perspectives in terms of extraction of valuable products from natural sources. Organic acids, for instance, are produced in large volumes for their use in biochemical and pharmaceutical industries. They are obtained as by-products of fermentation processes andhave to be separated from excess reagents and impurities. Traditional methods, such as extractive fermentation and selective precipitation, suffer from some drawbacks due to solvent toxicity, incompatibility in fermentation media and overall recovery costs. On the contrary, the use of SLMs offers many advantages in terms of lower energy consumption, higher selectivity, faster separation rate, less-toxic solvent and minimum backmixing effects [117]. The extraction of organic acids from kiwifruit juice by using a SLM process was investigated by Sch€afer and Hossain [118] as an alternative to the use of ion-exchange resins, electrodialysis and biological methods. The SLM consisted of an organic solution composed of a carrier (Aliquat 336/Alamine 336) and oleyl alcohol loaded on microporous polypropylene supports (Celgard 2500 and 2400). NaCl added to phosphate buffer solution to obtain a 1 M NaCl strip solution was used as the strip solution. Centrifuged and microﬁltered kiwifruit juice was used as feed solution. The permeability of the SLM system was evaluated in a batch cell, while the effect of

8.6 Conclusions

various process parameters, such as ﬂow rate of feed and strip solutions, carrier concentration and recycling mode of operation on the recovery of organic acids (quinic, citric, L-malic, L-ascorbic and fumaric) was analyzed in a continuous spiralwound membrane module. The results indicated that the ﬂux rates of organic acids decreased with the ﬂow rate of feed and strip solution and increased with the carrier concentration. Recycling of feed and strip solutions also resulted in a signiﬁcant improvement of the extraction efﬁciency. Citric, malic and quinic acids were extracted from kiwifruit juice at a rate of 5% in a single-pass process. One of the problems of the citrus industry is the formation of bitterness in citrus juice within hours after the extraction of the juice from the fruit due to the formation of bittering agents; limonoids and ﬂavonoids. Different methods including preharvest treatment with auxin plant-growth regulators, able to inhibit the biosynthesis of limonoids, extraction of limonoids and ﬂavonoids with cross-linked resin monomers and adsorption of bittering agent and acids with lignin-type adsorbents, have been developed to reduce or extract bittering agents from citrus juices. However, these methods present disadvantages that markedly limit their use. For instance, auxin plant-growth regulators are expensive and the adsoption of acid compounds along with bittering agents negatively affects ﬂavor. An extraction method of bittering agents in citrus juice, based on the use of membrane contactors, was proposed by van Eikeren and Brose [119]. In particular, the proposed approach was based on the use of two systems (A and B) of cellulose hollow-ﬁber membranes permeable to bittering agents and impermeable to ﬂavor and nutritional compounds. The citrus juice and an organic extractant were circulated in the shell side and in the lumen side of the system A, respectively. Bittering agents diffused across the membrane and dissolved into the organic extractant. The organic solvent enriched in bittering agents was transported and circulated into the lumen of the system B while a strip solution was circulated on the shell side of the same membranes. The bittering agents diffused across the membrane, and, being ionized by the basic strip solution, were trapped on the shell side. In these conditions a strip solution enriched in bitter agents and an organic solvent depleted in bitter agents (recycled to the lumens of hollow ﬁbers of the system A) were produced. The proposed system reduced the limonin concentration below the 6-ppm value, while the concentration of vitamin C remained unchanged. In another approach a supported-liquid membrane comprising a mixture of aliphatic solvents supported in the pores of a ﬂatsheet polypropylene membrane was used. The membrane was clamped between the feed solution and the strip compartment ﬁlled with a 0.01 M sodium hydroxide solution. The proposed system reduced the limonin content in the feed of 80%.

8.6 Conclusions

Table 8.5 summarizes main advantages and drawbacks of fruit-juice-concentration technique.

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Table 8.5 Advantages and drawbacks of fruit-juice-concentration techniques.

Process

Advantages

Disadvantages

Thermal evaporation

high evaporation ﬂuxes; high TSS content achievable; broad industrial scale application; possibility to use the same plant for different types of juice; the capital investment/evaporation capacity decreases by increasing the capacity aroma retention; no chemical alteration of juice components; organoleptic and nutritional preservation

high energy consumption; loss of organoleptic and nutritional properties; loss of aroma compounds; complex operation

Cryoconcentration

Reverse osmosis

broad industrial-scale application; low temperatures; modularity; easy scale-up; combination with vacuum evaporation; already commercially available; energetically and economically convenient if compared with thermal evaporation

Direct osmosis

low temperatures; low pressures; no fouling problems; constant permeate ﬂux in time; high TSS content achievable; modularity; easy scaleup; possibility to treat solutions with high suspended solids content; possibility to use the same unit to concentrate different products; low energy consumption; low cost of membrane replacement lower operating temperature with respect to thermal evaporation; reduced inﬂuence of concentration polarization in comparison to pressure-driven processes; theoretical 100% rejection to nonvolatile solutes; low energy consumption; high product quality lower operating temperature with respect to thermal evaporation; reduced inﬂuence of concentration polarization in comparison to pressure-driven processes; theoretical 100% rejection to nonvolatile solutes; low energy consumption; high product quality

Membrane distillation

Osmotic distillation

limit of concentration (30–50 Brix); necessity of an inactivation enzyme pretreatment; high investment costs; high energy consumption fouling phenomena; high pressures; necessity of an inactivation enzyme pretreatment; juice concentration limited at 22–23 Brix; loss of aroma compounds during the process; difﬁculty to concentrate solutions with high suspended solids content; high cost of membrane replacement new technology requiring an evaluation at industrial level; relatively low permeation (1.8–2.5 l m2 h1); high investment costs; possibility to use the same unit to concentrate different products

temperature polarization phenomena; heat losses by conduction through the polymeric membrane; lower transmembrane ﬂuxes in comparison with thermal evaporation

lower transmembrane ﬂuxes in comparison with thermal evaporation; brine disposal

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The potential advantages of OD and MD over conventional evaporation and other techniques have been successfully demonstrated at the lab scale and pilot scale, including improved product quality, lower energy consumption, and easy scale-up. With the enlargement of the worlds fruit-juice market and the request of product quality, commercial applications of membrane operations in fruit-juice processing, will expand in the near future. Within membrane operations membrane contactors can be considered applications in a breakthrough status with enhanced effectiveness. Their integration with standard membrane operations is a valid approach for a sustainable industrial growth within the process intensiﬁcation strategy. However, in order to gain a foothold in fruit-juice processing, development of suitable membranes with improved diffusional characteristics, selectivity, pore geometry and stability needs to be undertaken at reasonable costs.

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M.I. (1999) Temperature and concentration polarization in membrane distillation of aqueous salt solutions. J. Membrane Sci., 156, 265–273. Drioli, E., Jiao, B., and Calabrò, V. (1992) The preliminary study on the concentration of orange juice by membrane distillation. Proc. of VII International Citrus Congress, Acireale (Italy), 3, pp. 1140–1144. Cath, T.Y., Adama, V.D., and Childress, A.E. (2004) Expereimental study of distillation using direct contact membrane distillation: a new approach to ﬂux enhancement. J. Membrane Sci., 228, 5–16. Franken, A.C.M., Mulder, M.H.V., and Smolders, C.A. (1990) Pervaporation process using a thermal gradient as the driving force. J. Membrane Sci., 53, 127–141. Banat, F.A. and Simandl, J. (1994) Theoretical and experimental study in membrane distillation. Desalination, 33, 201–226. Phattaranawik, J., Jiraratananon, R., and Fane, A.G. (2003) Heat transport and membrane distillation coefﬁcients in direct contact membrane distillation. J. Membrane Sci., 150, 99–110. Lagana, F., Barbieri, G., and Drioli, E. (2000) Direct contact membrane distillation: modelling and concentration experiments. J. Membrane Sci., 166, 1–11. Gunko, S., Verbych, S., Bryk, M., and Hilal, N. (2006) Concentration of apple juice using direct contact membrane distillation. Desalination, 190, 117–124. Bagger-Jørgensen, R., Meyer, A.S., Varming, C., and Jonsson, G. (2004) Recovery of volatile aroma compounds from black currant juice by vacuum membrane distillation. J. Food Eng., 64, 23–31. Kozak, A., Bekassy-Molnar, E., and Vatai, G. (2009) Production of blackcurrant juice concentrate by using membrane distillation. Desalination, 241, 309–314. Calabrò, V., Jiao, B., and Drioli, E. (1994) Theoretical and experimental study on membrane distillation in the

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concentration of orange juice. Ind. Eng. Chem. Res., 33, 1803–1808. Diban, N., Voinea, O.C., Urtiaga, A., and Ortiz, I. (2009) Vacuum membrane distillation on the main pear aroma compound: experimental study and mass transfer modelling. J. Membrane Sci., 326, 64–75. Bandini, S. and Sarti, G.C. (2002) Concentration of must through vacuum membrane distillation. Desalination, 149, 253–259. Curcio, E., Barbieri, G., and Drioli, E. (2000) Direct contact membrane distillation in fruit-juice concentration. Ind. Bevande Italy, XXI, 113–121. Sarti, G.C., Gostoli, C., and Bandini, S. (1993) Extraction of organic components from aqueous streams by vacuum membrane distillation. J. Membrane Sci., 80, 21–33. Wang, Z., Zheng, F., Wu, Y., and Wang, S. (2001) Membrane osmotic distillation and its mathematical simulation. Desalination, 139, 423–428. Godino, M.P., Peña, L., Ortiz de Zarate, J.M., and Mengual, J.I. (1995) Coupled phenomena membrane distillation and osmotic distillation through a porous hydrophobic membrane. Separ. Sci. Technol., 30, 993–1011. Belaﬁ-Bakó, K. and Koroknai, B. (2006) Enhanced water ﬂux in fruit-juice concentration: coupled operation of osmotic evaporation and membrane distillation. J. Membrane Sci., 269, 187–193. Del Cero, C. and Boey, D. (1988) Liquid membrane extraction. Chem. Ind., 11, 681–687. Chimuka, L., Megersa, N., Norberg, J., Mathiasson, L., and J€ onsson, J.A. (1998) Incomplete trapping in supported-liquid membrane extraction with a stagnant

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acceptor phase. Anal. Chem., 70, 3906–3911. J€ onsson, J.A. and Mathiasson, L. (1999) Membrane extraction techniques in bioanalysis. Chromatographia, 52, S8–S11. Kemperman, A.J.B., Bargeman, D., Boomgaard, T.V.D., and Strathmann, H. (1996) Stability of supported-liquid membranes: state of the art. Separ. Sci. Technol., 31, 2733–2762. Alhusseini, A. and Ajbar, A. (2000) Mass transfer in supported-liquid membranes: a rigorous model. Math. Comput. Model., 32, 465–480. Mohapatra, P.K., Lakshmi, D.S., Mohan, D., and Manchanda, V.K. (2004) Selective transport of cesium using a supportedliquid membrane containing di-tbutylbenzo 18 crown 6 as the carrier. J. Membrane Sci., 232, 133–139. Di Luccio, M., Smith, B.D., Kida, T., Borges, C.P., and Alves, T.L.M. (2000) Separation of fructose from a mixture of sugars using supported-liquid membranes. J. Membrane Sci., 174, 217–224. Duan, S., Ito, A., and Ohkawa, A. (2003) Separation of propylene/propane mixture by a supported liquid membrane containing triethylene glycol and a silver salt. J. Membrane Sci., 215, 53–60. Crespo, J.G. and B€ oddeker, K.W. (1993) Membrane Processes in Separation and Puriﬁcation, NATO ASI Series, Kluwer Academic Publishers, Dordrecht, The Netherlands. Sch€afer, A. and Hossain, Md.M. (1996) Extraction of organic acids from kiwifruit juice using a supported liquid membrane process. Bioprocess Eng., 16, 25–33. van Eikeren, P. and Brose, D.J. (1993) Membrane extraction of citrus bittering agents, US Patent 5,263,409

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9 Membrane Bioreactors in Functional Food Ingredients Production Rosalinda Mazzei, Sudip Chakraborty, Enrico Drioli, and Lidietta Giorno 9.1 Introduction

The present chapter focuses on the application of membrane reactors using catalysts of biological origin for food productions. An overview about the different membrane bioreactor types is reported, and their advantages together with the main drawbacks are discussed. The use of membrane bioreactors in the different food applications is described with more attention in the production of functional food.

9.2 Membrane Bioreactors and Functional Food

During food processing, ﬂavor and odor are often lost, obtaining as ﬁnal products lower food quality if compared with fresh ingredients. The development of better methods for delivering ﬂavor is of high interest for the food industry to heighten user enjoyment, this is particular important if we consider continuing urbanization and increasing problems in transportation due to energy consumption. An important sector that contributes in the food industry to this aim are separation techniques that can isolate ﬂavor and odor chemicals early in the processing steps and resupply them to the processed foods. Functional food processes can recover small components active ingredients from by-product streams to be used as high-value additives. Membrane bioreactors are able to integrate bioconversions with selective membrane separations leading to continuous clean, safe and low energy consumption production. Although their potentialities have not been fulﬁlled yet, the current needs and challenges in satisfying the increasing consumer demand of safe goods and the limited resources availability will force the industry towards these selective and efﬁcient techniques. The use of biocatalysts in combination with membrane operations permits drawbacks to be overcome enabling biotransformation to be integrated into contin-

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uous production lines. These systems, being able to work at time-invariant conditions at steady state, permit a better control of reaction conditions with an increase of lifetime, productivity and economic viability of the process. In addition, the separation, puriﬁcation and concentration of the obtained product can be obtained. Thanks to the biocatalyst and membrane selectivity the mass intensity can be very high, with no by-products formation, while producing high added value coproducts. Membrane processes and in particular membrane bioreactors are regarded as particularly suitable for food applications because, in general they can operate under mild conditions of temperature, pressure and shear stress, therefore preserving the biological activity of the compounds to be recovered and the properties of the original media/matrix. In general, they do not require any extraction mass agent such as solvents, avoiding product contamination and the need for subsequent puriﬁcation. Based on the membrane role, bioreactors are divided into systems in which the membrane does not contribute to the reaction but only controls mass transport, and systems in which the membrane works as a catalytic/separation unit, a conﬁguration in which the reaction also occurs at the membrane level. In this last case we talk about biocatalytic membrane reactors, BMR. The ﬁrst case represents the most commonly used, and due to the presence of different biocatalysts (enzyme or cells) of different molecular weight, in the literature several names are found to describe this. In this work, to indicate membrane bioreactor in which the membrane acts as separation unit, we will refer to free biocatalysts membrane bioreactors (MBR), in the other case we will refer to BMR. A schematic representation is reported in Figure 9.1. In biocatalytic membrane reactors the biocatalyst can be: entrapped, geliﬁed, and bound to the membrane. Biocatalytic membrane reactors with biocatalyst bound to the membrane can result from ionic binding, cross-linking and covalent binding. Membranes in a variety of conﬁgurations, including tubular, hollow ﬁber and spiral wound have been used in the food industry for many years. They can be applied within the production process, that is for clariﬁcation and concentration, as well as to treat the resulting wastewater prior to disposal or reuse. The main beneﬁts of membrane technology/bioreactor are well documented. Examples of systems Membrane Bioreactors MBR

BMR

Methods to heterogenize enzyme on membrane

Entrapment

Figure 9.1 Schematic representation of membrane bioreactors.

Gelification

Binding to the membrane

9.3 Membrane Bioreactor in Sugar and Starch Processing

developed at industrial production level have been recently reviewed [1]. However, more research effort needs to be invested in order to fully exploit their potentialities. Their implementation falls primarily in two main objectives (a) improve production process (b) recovery of valuable products that previously would have been lost as wastes. A new trend in the development of membrane bioreactor is dictated from the strong need in food/feed to produce functional food. Functional food or medicinal food is any fresh or processed food claimed to have a health-promoting or disease-preventing property beyond the basic function of supplying nutrients. To better explain the contribution of membrane bioreactors development in functional food production some deﬁnitions are outlined. A nutraceutical is a part of functional food isolated and puriﬁed from foods that has physiological beneﬁt or provides protection against chronic diseases. Bioactive compounds are examples of nutraceuticals. Treated food with live cultures are considered as functional food with probiotic components, which is a viable microbial dietary supplement, that beneﬁcially affects the host through its effects in the intestinal tract. In other words, probiotic foods are deﬁned as those that contain a single or mixed culture of micro-organisms that affect beneﬁcially the consumers health by improving their intestinal microbial balance [2]. Another class of treated food is prebiotic. The term was ﬁrst used in Japan in the 1980s where there is a government approval process for functional foods called Foods for Speciﬁed Health Use (FOSHU). A prebiotic is a food ingredient that is not hydrolyzed by the human digestive enzymes in the upper gastrointestinal tract and beneﬁcially affects the host by selectively stimulating the growth and/or activity of one or a limited number of bacteria in the colon that can improve host health [3]. The ﬁbers are included in this kind of compounds.

9.3 Membrane Bioreactor in Sugar and Starch Processing

The sugar and starch industries represent the competitive world sugar market, Table 9.1 summarizes major sugar productive countries and their production. Application of membrane technology in the sugar industry contributes to the sustainable development in the ﬁeld. In particular, clariﬁcation of sugarcane juice, production of glucose or glycerol as well as are sugar-related products are aspects where membrane bioreactor can play a role. In the confectionery and many food and beverage industries, sugar (present as starch, sucrose, fructose and glucose, etc.) is the main constituent in some of the process streams. Inevitably, it is also present in the efﬂuent streams arising from these industries. There is, however, considerable interest among manufacturers to optimize process economics through product recovery, and to respond to environmental pressure to reduce the waste generation. Enzymatic hydrolysis of starch is traditionally performed in large volume-batch reactors using soluble enzyme following a two-step procedure including the

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Table 9.1 Major sugar-producing countries.

2007/08 est.

Production (million tons)

Exports (million tons)

Per capita consumption (kg)

Brazil India EU China Thailand United States Mexico SADC Australia Pakistan

31.355 28.804 17.567 14.674 8.033 7.701 5.978 5.834 5.013 4.891

20.957 (1) 3.298 (4) 1.400 (8) — 5.288 (2) — 0.350 (15) 2.410 (5) 3.750 (3) —

58 20 34 11 36 29 52 22 47 25

liquefaction and sacchariﬁcation [4]. But nowadays it can be performed in a single step by using amylase enzyme termamyl that is able to produce dextrins [5, 6] using membrane reactor. First, the starch being dissolved in water and partially hydrolyzed with an a-amylase to give maltodextrines and in the next step, sacchariﬁcation enzymes transform liqueﬁed starch into low molecular weight oligosaccharides such as glucose or maltose. The conventional batch reaction processes has a great number of disadvantages, such as incompatibility of enzyme recovery and reuse, high labor and puriﬁcation cost, high capital investment and discrepancies in glucose syrup quality, low efﬁciency, batch to batch variations, and most of all, the high enzyme cost. But the application of membrane reactors make possible continuous operation in lower reactor volume, as well as in shorter reaction time [5–9] an increase the reactors efﬁciency and ﬁnally to reuse enzymes in a continuous way. The application of membrane reactors in the starch industry is particularly used in the production of smaller assimilable sugars. This reaction is carried out in systems in which the enzyme is not immobilized and the membrane works as a separation device (MBR). The enzymes used are amylolytic enzymes and debranching enzymes [10] or using liqueﬁed starch as substrate [11]. The major problems in the application of the membrane reactor are a large decrease in permeate ﬂux due to concentration polarization and fouling [5, 6, 12]. Different solutions are applied to decrease fouling phenomena such as the pretreatment of the raw starch solution. Other studies were devoted to the examination of factors that mainly affect membrane performance such as: molecular-weight cut-off, enzyme dosage, residence time, transmembrane pressure, carbohydrate composition, and retention factor [11]. Membrane bioreactors have been used for production of glucose, maltose, maltotetraose, and cyclodextrins [5–9, 13–16] for the food-grade industrial production [12] like puddings, jellies, and fruit desserts. In this system hydrolysis can be carried out simultaneously by separating syrups from enzymes and non-hydrolyzed starch [12]. An extra separator system to extract the product is not necessary but it is needed to

9.3 Membrane Bioreactor in Sugar and Starch Processing Table 9.2 Examples on the use of membrane bioreactors for sugar production.

Starch used

Enzyme used

Membrane used

Commercial potato starch Cassava Cassava

a-amylase (BAN 480L) Termamyl Maltogenase and promozyme Amylolitic enzyme complex – from fermentation of whole wheat ﬂour by Aspergillus awamori

Tubular ceramic

Amylos

Molecular weight cut-off (kDa)

Reactor type

Reference

50

MBR

[17]

Carbosep M4 Carbosep M4

50 50

BMR BMR

[18] [19]

Hydrophilic cellulose acetate

40

MBR/ BMR

[20]

concentrate the product by application of different advanced membrane operations (UF/NF). Table 9.2 summarizes examples of membrane bioreactors used for sugar production. Another ﬁeld of membrane bioreactor application is the production of cyclodextrins or oligosaccarides. The development in this ﬁeld was pushed from the high interest devoted to this compound in the last period, due to the fact that they have applications in several ﬁelds, including food, pharmaceutical, cosmetic, and plastic industries as emulsiﬁers, antioxidants and stabilizing agents. The production of cyclodextrins by membrane bioreactors was conducted using different starting sources including corn starch and soluble potato. A recent work reported their production also starting from tapioca starch [21]. The production of oligosaccarides to be used as functional food was also obtained by the immobilization of dextranase on polymeric matrix [22]. Cyclodextrins can be used as carriers for molecular encapsulation of ﬂavors and other sensitive ingredients [23]. The molecular encapsulation of lypophilic food Table 9.3 Examples of laccase immobilization on different membrane material.

Biocatalyst

Membrane material

Immobilization

Reference

Laccase from Aspergillus sp. Laccase from Pyricularia oryzae Laccase from Trametes versicolor Laccase from Trametes versicolor

Nylon-66 Polyethersulfone membranes Hydrophilic PVDF microﬁltration membrane Polyether sulfone membranes

Adsorption Adsorption

[42, 44] [45]

Covalent binding

[46]

Entrapment

[47]

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ingredients with cyclodextrin improves the stability of ﬂavors, vitamins, colorants and unsaturated fats, and so on. Various types of oligosaccharides have been found as natural components in many common foods including fruits, vegetables, milk, honey. Oligosaccharides can also be used as functional food ingredients that have a great potential to improve the quality of many foods. In addition to providing useful modiﬁcations to physicochemical properties of foods – such as the improvement of intestinal microﬂora based on the selective proliferation of biﬁdobacteria, stimulation of mineral absorption, non- or anticariogenicity, and the improvement of both plasma cholesterol and blood-glucose level. Basically oligosaccharides are short-chain sugars generally consisting of two to ten building block small sugars. It is used as a nutrition supplement in food ingredients and additives. Apart from direct extraction from plants the oligosaccharides can be processed by enzymatic synthesis using enzymes that possess hydrolytic or transglycosylation activity, in continuous membrane bioreactors. Both batch reactor with soluble enzymes and continuous systems with enzymes or whole cells immobilized have been used.

9.4 Membrane Bioreactor in Oil and Fat Processing Industry

The use of membrane bioreactors for the hydrolysis of oils and fats is intensively investigated. The biocatalysts used are mainly lipases and esterases and the processes in which they are involved for functional food production are ester synthesis to produce emulsiﬁers and aroma compounds and oil hydrolysis for free fatty acids, mono or diglycerides productions. Monoglycerides, diglycerides, triglycerides, and glycerol are widely used in the food industry as emulsiﬁers for bakery products, margarines, dairy products, confectionery, and so on. In foods and beverages, glycerol serves as a humectant, solvent and sweetener, and may help preserve foods. It is also used as ﬁller in commercially prepared low-fat foods (e.g., cookies), and as a thickening agent in liqueurs, although it has about the same food energy as table sugar. Glycerin has many uses, such as in the manufacture of food and in the production of pharmaceuticals too. The most commonly used products are glycerol monostearate, monooleate, and monoricinoleate [24]. The complex mixtures that contain 40–48% monoglycerides (MG), 30–40% diglycerides (DG), 5–10% triglycerides (TG), 0.2–9% free fatty acids (FFA), and 4–8% free glycerol are generally termed monoglycerides. These mixtures have applications in food fats (margarine, ice cream, sweets, etc.). Pure monoglycerides (90–97%), obtained by molecular distillation of the above mixtures, are also commercially available. The higher-purity monoglycerides are preferred for bakery uses because of their good amylase complexing ability. Most commercial MG are produced from edible, reﬁned, hydrogenated animal fats (tallow, lard, etc.) or from hydrogenated vegetable oils (palm, soybean, corn, olive, peanut, etc.). High oleic vegetable oils can also be used as raw materials for the production of emulsiﬁers for liquid and low-fat margarines.

9.4 Membrane Bioreactor in Oil and Fat Processing Industry

The monoglycerides can be produced on an industrial scale by glycerolysis of fats and oils by means of inorganic alkaline catalysts, such as sodium hydroxide or by enzymatic routes. Application of enzymes as catalysts for reactions in the oils and fats industry is being extensively studied in the literature. Enzymes are chosen since they show many advantages over traditional inorganic catalysts: they have large catalytic activity under mild operating conditions; they show large selectivity to the desired product with no signiﬁcant side reactions, leading to products of high purity. Various review papers about membrane bioreactors using lipase in vegetable oil and fat processing have been published during recent decades [25–27]. The inﬂuence of operating conditions of lipase immobilized in a two-separate phase membrane bioreactor has been reported [29–31]. In particular, the effect of immobilization method, amount of enzyme, hydrodynamic conditions, and microenvironment conditions (such as pH, temperature, membrane material) have been investigated [32]. Strategies to improve reaction performance as well as transport properties through the enzyme-loaded multiphase system have been exploited [33, 34]. Three different enzyme membrane reactors have been compared, as illustrated in the Figure 9.2. Lipase was used free in a stirred-tank reactor, and as immobilized in

Figure 9.2 Schematic representation of bioreactors studied [29–31]using lipase as biocatalyst: free stirred-tank reactor, two-separate phase enzyme membrane reactor and two-separate phase enzyme membrane reactor with emulsions.

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a membrane in the absence and in the presence of oil/water droplets. The use of oil droplets immobilized together with the enzyme signiﬁcantly improved the performance of the system thanks to the positive effect of the o/w interface uniformly distributed through the membrane on the enzyme activity as well as on the substrate transport. Lipase has been immobilized on polymer membranes with hydrophilic [35] hydrophobic [36] properties, as well as on inorganic membranes [37]. Another application of membrane bioreactors is production of speciﬁc structured lipids in an enzymatic route from rapeseed oil and capric acid [38]. Production of v3polyunsaturated fatty acid (v3-PUFA) concentrates from ﬁsh liver oils (which have been claimed to provide beneﬁcial health effects via prevention of coronary heart diseases) for use as nutropharmaceutical food supplements is another application of lipase in membrane bioreactor, and sequential lipase-catalyzed chemical incorporation in triglycerides. Lipase from Candida rugosa was also immobilized on Cuprophane membrane [39] in a hollow-ﬁber module.

9.5 Membrane Bioreactors in Hard Drink Industry and Liquid Beverages 9.5.1 Wine

Membrane bioreactors are developing in the wine ﬁeld for the production of aromatic compounds and ﬂavor by the use of glucosidases, the production of additives from pectinase hydrolysis, and the production of preservatives molecule such as lactic acid by the use of malolactic bacteria. The production of wine in terms of cropped surfaces and product yield ﬂuctuates in a signiﬁcant way over the years. At the end of the 1990s the production tended to decrease, but a signiﬁcant increase was achieved at the end of 2004/2005 going back toward another decrease in 2005/2006, where the production was 4% less. Nevertheless, the production of European countries (27 countries) in 2007 was about 174 449.170 (Wine, production – 1000 hl). With respect to the total worldwide production, Europe represents the higher producer of wine (http://news.reseauconcept.net/images/oiv/client/STATISTIQUE__Verone_2008_EN_deﬁnitif_41diapos.pps#1) having leading countries like France, Italy and Spain. The United States is another important producer followed by Argentina and China, while the economy of other countries like Germany, South Africa and Chile, is growing in the last three years. Thanks to the action of different yeasts, both Saccaromyces and non-Saccaromyces type, in the ﬁrst part of wine making there is the conversion of glucose in ethanol CO2 and other products. The presence of the yeast is fundamental in must fermentation due to the production of particular enzymes that help the fermentation process. The use of these type of enzymes or directly the yeasts with this enzymatic activity, immobilized with membrane or on other support for wine fermentation is of high interest.

9.5 Membrane Bioreactors in Hard Drink Industry and Liquid Beverages

Some examples of coupling the enzymes useful in wine making and membrane reactors are reported in Table 9.4. b-glucosidase is an important enzyme in wine making, the enzyme is employed in different applications like production of rose wine from red grapes, for the hydrolysis of antocianines, and for the hydrolysis of terpenglucosides and so on [40]. The immobilization of this enzyme, or bacteria and yeast showing that enzymatic activity is of high interest in beverages production with enhanced aroma. In the literature are reported some examples [40] about the immobilization of b-glucosidase on different support (Cellulose PEI (Baker), alpha-aluminia CT 2000 (Alcoa Chemie), gamma-aluminia (Akzo), chitosan (Chitobios) and polymeric) applying adsorption, covalent bonding by glutaraldeide and cross-linking immobilization techniques. Some immobilized glucosidase enzymes has also been proved on a pilot scale. They were used on a continuous-ﬂow stirred-tank membrane reactor in a model system and also during wine making [41]. In this system the enzyme was immobilized on chitosan pellets and to simulate the natural process, the medium was also supplemented with chemicals present in the wines (fructose, ethanol, nerol, linanol, geraniol). Fructose did not decrease biocatalyst stability, while alcohol affected the enzyme half-life from 2586 h at 3% (w: v) ethanol to 1378 h at 12% (w: v). Enzyme stability was not dependent on substrate concentration and was considered satisfactory for an industrial process (a half-life of 1.2 years). Many precursors of the aromatic components of wine are monoterpenes (geraniol, nerol, citronellol, linalool, a-terpineol, etc.) in di-glycosidic form, that contain b-Dglucopyranose bound directly to aglycon and/or other sugars among which are a-Lrhamnopyranose and a-L-arabinofuranose. Therefore, to develop the aromatic potential of a wine to the full, together to ramnopyranose (Rha), it is also necessary to utilize the other glycosidases: a-Larabinofuranosidase (Ara, EC 3.2.1.55), and ﬁrst b-D-glucopyranosidase (bG, EC 3.2.1.21) and a-L-rhamnopyranosidase (Rha, EC 3.2.1.40). An important reaction that occurs in wines and in particular in white wine and in rose is the development of madeirized ﬂavor. The process is mainly due to polyphenols, that can have also beneﬁcial health effect because of their antioxidant properties. Oxidative enzymes like laccase coming from fungi are used to improve the process. Several studies were performed on the use of Laccase in phenol-removal processes for must and wine stabilization [42]. Laccase was immobilized on different membrane materials by applying different immobilization techniques from different sources (Table 9.3). Cantarelli and Giovanelli [43] carried out assays in order to determine if the enzymatic preparations could be used in white-wine production for polyphenols reduction in musts (and consequent stabilization of the wine color) instead of oxidation. The results demonstrated that the enzymatic treatment coupled with ﬁltration with polyvinylpolypyrrolodone (PVPP) reduced the quantity of oxidized polyphenols. Other important enzymatic activities in wine making and in particular in wine-clariﬁcation processes are pectinases, which are usually used to improve

j209

Examples of membrane bioreactors for pectins hydrolysis.

Biocatalyst

Membrane

Bioreactor configuration

Application

Reference

Pectin lyase from Penicillum italicum

Ultraﬁltration membrane

MBR

[48]

Polygalacturonase from A. niger

30-kDa ﬂat regenerated cellulose membrane Spiral-wound polysulfone membrane (10 kDa) Amicon 10 kDa

MBR MBR

Production of pectic oligosaccarides Production of D-Galacturonic acid Wine clariﬁcation

Titania microﬁltration

BMR

Polyvinilidene ﬂuoride tubular, polysulfone spiral wound Polysulfone single-hollow ﬁber Hollow-ﬁber ultraﬁltration 10-kDa spiral-wound polysulfone

Polygalacturonase and pectin lyase from A. niger Endo-polygalacturonase from Aspergillus pulverulentus Polygalacturonase from A. niger Rapidase liquid plus Amylase and pectinase Commercial pectinaase Endopectidase from A. niger

MBR

[49] [50] [51]

BMR

Production of pectic oligosaccarides Production of pectic oligosaccarides Apple-juice clariﬁcation

BMR

Fruit-juice processing

[54]

BMR MBR

Fruit-juice processing Apple-pectin hydrolysis

[55] [50]

[52] [53]

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Table 9.4

9.5 Membrane Bioreactors in Hard Drink Industry and Liquid Beverages

processability and to produce additives. The main membrane bioreactor conﬁguration used in the wine industry using pectinases action is a free enzyme membrane reactor (BMR). The soluble enzyme is conﬁned in the retentate side of the membranes where it is in contact with the substrate. In Table 9.4 some examples and membrane material used for the pectines hydrolysis are reported. These applications are referred both to the wine and fruit-juice treatment. Together with protein immobilization, the alternative strategy for wine making, is cell immobilization. Although this application is rapidly expanding in the research area, development at the industrial scale is still limited. Takaya et al. [56] studied the efﬁciency of two membrane bioreactor systems for continuous dry wine making. The ﬁrst conﬁguration was a single-vessel bioreactor, while the second conﬁguration included two vessels; one operated as a continuous stirred-tank reactor and the other was a membrane bioreactor. The double vessel resulted in 28 times more productive that the single one. Cell immobilization is a rapidly expanding research area. The purpose of this technique is to improve alcohol production and overall product aroma, taste and quality. Many support are used for cell immobilization in this ﬁeld divided into inorganic, organic and natural materials. Some examples of different supports and its main application are reported in Table 9.5. Malolactic fermentation is a secondary process that occurs in wines during the maturation period. Lactic bacteria predominately of the genera Oenococcus, Lactobacillus and Pediococcus are responsible of this process, where L-malic acid is converted to lactic acid, an important food preservative, and carbon dioxide. As a consequence of this reaction the total acidity of the wine decreases. Oenococcus oeni can carry out this process in one step, without the production of piruvic acid. Other by-product produced during this fermentation can affect wine ﬂavor. Also some yeast as Saccaromyces can convert malic acid through maloethanolic fermentation [65]. The immobilization technology is important also in this ﬁeld, where the cell compartimentalization can help to (i) increase the tolerance towards malolactic

Table 9.5 Materials used for cell immobilization.

Inorganic material Mineral kissis c-aluminia Organic support Cellulose covered with Ca-alginate Ca-alginate beds Natural support Deligniﬁed cellulose Gluten pellets

Immobilized cell

Application

Reference

Saccaromices Saccaromices

Aroma improvement Aroma improvement

[57] [58]

Saccaromices and Candida Saccaromices

Enhance glycerol formation in wine Must fermentation

[59, 60]

Saccaromices

Fermentation Production of wine with less alcohol content

[62–64]

[61]

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fermentation bacteria, (ii) develop the desired ﬂavor selecting the appropriate cultures, (iii) accelerate the process increasing cell densities, (iv) reuse of the cell. A kinetic analysis was carried out using three different immobilization techniques of malic enzyme for the development of a membrane bioreactor: (1) polymeric membranes [66] and cross-linking reaction, (2) within polyurethane foams, and within a gel-like membrane formed on active side of ultraﬁltration polymeric membranes [67]. Enzymatic cell-free reactors, did not allow to efﬁcient, complete and rapid consumption of the L-malic acid to be achieved [68, 69]. 9.5.2 Beer

Beer is the second most consumed beverage in the world behind tea, and it continues to be a popular drink. The brewing industry has an ancient tradition and is still a dynamic sector open to modern technology and scientiﬁc progress. Brewers are very concerned that the ﬁnishing techniques they use are the best in terms of product quality and cost effectiveness [70]. Beer production requires about seven days of fermentation and large-scale fermentation and storage capacity. The main ﬁeld in which membrane bioreactors can be developed in beer are the alchohol-free beers and in the maturation and aroma control. In the ﬁrst process, the two main approaches currently used are the removal of the alcohol from product and limited fermentation. In the case of limited fermentation the system is most efﬁcient where the fermentation cells are immobilized. The yeasts commonly used for this process are S. cerevisiae. Different kinds of support are used to immobilize the yeasts in brewing, they can be divided in inorganic, organic and natural. The prevalent organic support are: polyethylene, PVC, polysaccarides, DEAE-cellulose; the inorganic porous ceramic and silicon and the natural support are deligniﬁed cellulose and gluten pellets [71]. 9.5.3 Ethanol Production

The requirement of ethanol in the beverage industries as an additive has been steadily increasing and so is the pursuit of immobilized microbial cell systems for ethanol fermentation. Research on alcohol production usually focuses on volatile by-product formation, because these constituents are critical parameters for distillates and alcoholic beverage quality. For ethanol production different yeast strains are used such as: S. cerevisiae, S. diastaticus, K. marxianus and Candida sp., and different bacteria like Zymomonas mobilis. The requirement for food-grade purity is not essential due to the employment of a distillation step. A membrane distillation bioreactor was developed for ethanol production [72, 73], where the batch fermentation was coupled with a membrane distillation process. The porous capillary polypropylene membranes were used for the separation of volatile compounds from

9.6 Membrane Bioreactor in Other Liquid Beverages

the feed. The elimination of these compounds allows an increase in ethanol productivity and rate. In this case the yeast used was S. cerevisiae. A membrane bioreactor for the production of ethanol was developed in a pilot plant [74]. This system integrated ceramic microﬁltration membranes with a stirredtank bioreactor.

9.6 Membrane Bioreactor in Other Liquid Beverages

The main applications of membrane bioreactors in other drink industries are: reducing the viscosity of juices by hydrolyzing pectins, reducing the lactose content in milk and whey by its conversion into digestible sugar. 9.6.1 Fruit-Juices Production

The production of fruit juices is divided into six major steps: crushing, pressing, clariﬁcation, centrifugation or ﬁltration, concentration, pasteurization. During the fruit crushing there is the solubilization of pectins, these compounds can usually affect the processability, creating turbidity and cloud forming. Pectinase, the pectolitic enzyme responsible for pectins hydrolysis are commonly used in the fruit-juice industry, in two steps: pressing and clariﬁcations. During pectin hydrolysis the monomer of pectin, D-galacturonic acid is also produced, which is an important compound, as a raw material in the food, pharmaceutical and cosmetic industries to manufacture, for example vitamin C, or acidifying, tensioactive agents. Oligosaccarides derived from pectin hydrolysis can also have some important applications as repressors of liver lipid accumulation in rats [75], as antifungal phytoalexin-elicitors in plants [76], inducers of ﬂowering and antibacterial agents [115]. Traditionally, enzymatic hydrolysis of pectins has been conducted in batch systems. Unfortunately, after each cycle of operation the enzyme could not be recovered for further use and immobilized enzyme could suffer from steric hindrance effects and losses in enzyme activity as a result of immobilization. The use of membrane bioreactor is the alternative efﬁcient strategy, in which the enzyme is retained or compartmentalized, thus increasing enzyme utilization. One of the membrane bioreactor conﬁgurations commonly used is with the enzyme compartmentalized in the retentate side of the membrane together with the substrate, while the product is separated in the permeate. Different works were carried out for pectin hydrolysis in membrane bioreactor systems using a free enzyme membrane reactor. Alkorta et al. [77] studied the reduction in viscosity of pectins catalyzed from pectin lyase from Penicillum italicum in a membrane reactor. This enzyme results as the only pectinase enzyme capable of hydrolyzing a-1,4 glycosidic bond of highly esteriﬁed pectins, without altering the

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volatile compounds responsible for the aroma of various fruits [77, 78] the reduction in viscosity was demonstrated with high efﬁciency towards different fruit juices: grape, peach, melon, apple and pear, showing a little decrease in the case of apple and pear juice. Another biocatalyst used frequently in pectin hydrolysis was polygalactorunase from A. niger. A. niger pectinases are most widely used in industry because this strain possesses GRAS (generally regarded as safe) status, so the metabolites coming from its production can be directly used without further treatment [79]. The pectinases produced from this strain are: polymethylgalacturonase (PMG), polygalacturonase (PG) and pectinesterase. However, particular pectinases are used for speciﬁc purpose, for example only polygalacturonnase is used for baby-food products [79]. A recent work reports the use of polygalactorunase from A. niger in a ﬂat-sheet membrane reactor, which shows excellent stability for more than 50 h. In this case, the membrane used was a 30-kDa regenerated cellulose membrane [49]. The same biocatalyst was used in a free enzyme membrane reactor where the membrane used was a spiral-wound polysulfone membrane (10-kDa MWCO), attaining a conversion of 83% and a stability for a long period (15 day) [50]. The performance of pectin hydrolysis was also tested by immobilizing directly the enzyme on the membrane and conducting the reaction in a biocatalytic membrane reactor [77]. The use of pectinases immobilized on ultraﬁltration membrane hydrolyze the pectin to lower molecular weight species, permitting an extension of membrane operation without cleaning [55]. Pectinase was also immobilized by physical immobilization on a titania microﬁltration membrane [52] and on a polysulfone hollow-ﬁber membrane [55], and coimmobilized with amylase on a polymeric hollow-ﬁber membrane to hydrolyze simultaneously starch and pectins. The coimmobilization showed an improvement of ﬂux of an additional 35% [54]. An integrated membrane process for producing apple-juice and apple-juice aroma concentrates was proposed by Álvarez et al. [80]. The efﬁcient system involves the following operations: an integrated membrane reactor to clarify the raw juice; reverse osmosis to preconcentrate the juice, pervaporation to recover and concentrate the aroma compounds, and ﬁnal an evaporation step to concentrate apple juice. These operations were tested in laboratory and pilot-plant units, giving promising results both on the yield of product and also for economical aspects. Some examples of immobilized pectic enzyme are present at the industrial scale [29, 30, 48] 9.6.1.1 Functional Food Production in the Milk and Whey Field by Membrane Bioreactor The ﬁrst application on a large scale of a membrane bioreactor was the hydrolysis of lactose by immobilized b-galactosidase on a cellulose ﬁber for the production of milk with low lactose content [81]. Lactose, together with high molecular weight proteins, are allergenic compounds present in both milk and whey. Intolerance to milk comes from the fact that some subjects cannot digest proteins, contained in milk and whey, with a molecular weight higher than 5 kDa.

9.6 Membrane Bioreactor in Other Liquid Beverages

9.6.1.1.1 Lactose Hydrolysis Lactose is the dominant carbohydrate in milks and it is also contained in whey. A large number of people do not digest lactose properly due to the lack or inactivity of the intestinal b-galactosidase and they suffer from intestinal dysfunction. In addition, lactose is a sugar with a high BOD, low sweetness, and low solubility and has a strong tendency to adsorb ﬂavors and odors compared to its hydrolysis products; glucose and galactose. Lactose hydrolysis is an important food process, not only to produce lactose-free milk, but also to improve processes for the production of refrigerated diary products, because some technological difﬁculties occurs associated with lactose crystallization [82]. Another important application of lactose hydrolysis is the production of additives, like lactic acid, glucose and galactose that can be used in the human diet [83]. For the industrial applications of enzymes to the productions of large quantities of product, the enzymes should be immobilized to be used in continuous reactors. Several procedures for b-galactosidase have been studied: entrapment, adsorption, ionic interaction, afﬁnity, complex formation with metal, and covalent bonds [83]. Several reactors were also tested using different membrane reactors conﬁguration and different starting sources. In Table 9.6 some examples showing support material application are reported. The main enzyme used in membrane bioreactors for lactose hydrolysis are from Kluyveromyces yeast and Aspergillus fungi, micro-organisms considered safe (GRAS). In particular the enzymes from fungi can be used in acid wheys since their optimum pH is 3.5–4.5, while the enzymes from yeasts can be used in milk and sweet wheys since their optimum pH is between 6.5–7 [84]. As previously mentioned, the other application of membrane bioreactors in the lactose hydrolysis is the production of lactic acid. Lactic acid is one of the value-added product produced from processing cheese whey. The food and drug administration have approved lactic acid and its salts to be GRAS. The bacteria usually used for the production of lactic acid by fermentation process from cheese whey are Lactobacillus helveticus [91–93] and Lactobacillus casei, while Biﬁdobacterium longum converts lactose into lactic acid and produces antibacterial compounds [94]. The main conﬁguration of a membrane bioreactor for the production of lactic acid is a fermentation reactor with a membrane unit as reported in Table 9.7. in this kind of conﬁguration cell, protein and lactose are separated by a ﬁltration unit and returned to the fermentor while lactic acid is separated in the permeate. Some examples of biocatalytic membrane reactors are also present in the literature. L helveticus cell were immobilized in a polymeric membrane reaching a lactose conversion of 79% and a lactic acid yield of 0.84 g of lactic acid/g of lactose utilized [97]. A two-stage continuous fermentation with membrane recycle has been studied that enhances lactic acid productivity from 21.6 g dm 3 h1 in a single stage to 57 g/ dm3 h1 in two stages [95]. 9.6.1.1.2 Protein Hydrolysis in Milk and Whey by MBR The hydrolysis of high molecular weight proteins into small polypeptides is an alternative approach to produce low allergenic (b-lactoglobulin) fresh milk. The possibility to hydrolyze high molecular weight proteins by membrane bioreactors provides a rich source of peptides that are latent until released and

j215

Examples of membrane bioreactor used to hydrolyze lactose.

Biocatalyst

Source

Material and reactor configuration

Application

Reference

B. circulans K. lactis, A. oryzae b-glycosidases from the archaea Sulfolobus solfataricus (Ss b Gly) and Pyrococcus furiosus (CelB) b -galactosidase from Kluyveromyces lactis b-galactosidase commercial enzyme A. orza, K. lactis

Skimmed milk Lactose Lactose

MBR MBR with ceramic membrane MBR with an ultraﬁltration unit

High-quality milk Production of galactosil-oligosaccarides Production of oligosaccarides

[85] [86] [87]

Lactose

BMR

Galactose and glucose production

[88]

Lactose Lactose

MBR MBR

Production of oligosaccarides Production of Galactosyl-oligosaccharides

[89] [90]

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Table 9.6

9.6 Membrane Bioreactor in Other Liquid Beverages Table 9.7 Examples of membrane bioreactors in the production of lactic acid.

Biocatalyst

Source

Membrane-reactor configuration

Reference

L. ramnosus L helveticus L. casei L. ramnosus

Glucose Whey Lactose Lactose

MBR MBR MBR MBR

[96] [97, 98] [99] [97, 98]

activated, for example, during gastrointestinal digestion or milk fermentation. Once activated, these peptides are potential modulators of many regulatory process. Milk-derived bioactive peptides can have physiological functionality on cardiovascular, (antihypertensive, antioxidative, antithrombotic, hypocholesterolemic), nervous (agonistic, anthagonistic oppioid activity), gastrointestinal (antiappetizing, antimicrobial) and immune (antimicrobial, immunomodulatory, citomodulatory effect) systems [100]. The active peptides can be produced by the hydrolysis of digestive enzymes, through proteolytic micro-organism and through the action of proteolytic enzymes derived from micro-organisms or plant. Some examples are reported in Table 9.8. Commercial production of bioactive compounds from milk proteins is limited. The use of enzymatic membrane reactors for continuous production of speciﬁed peptide sequences was introduced during 1990. Nowadays, it has been widely studied, in the literature, for total conversion of food proteins of various origins with improved nutritional and/or functional properties. Continuous extraction of bioactive peptides in membrane reactors has been mainly applied to milk proteins using different membrane material and different membrane reactor conﬁguration (See Table 9.9).

Table 9.8 Examples of biocatalyst used to produce active peptides from protein source.

Biocatalyst

Protein source

Active peptides produced

Reference

Pepsin Trypsin

Casein Casein

[100] [101, 102]

Protease N Lactococcus lactis Lactococcus helveticus

Whey protein Casein, milk Casein, whey proteins Casein

(ACE) inhibitory peptides (ACE) inhibitory calcium-binding phosphopeptides Different peptides (ACE) inhibitory peptides (ACE) inhibitory peptides (ACE) inhibitory peptides

[105]

Lactobacillus delbruecki ssp. vulgaris

[103] [100] [104]

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Examples of production of bioactive peptides using MBR.

Biocatalyst

Substrate

Membrane reactor configuration

Application

Reference

Alcalase Trypsin

Casein Caseinomacropeptides

MBR MBR with ultraﬁltration unit

[106] [107]

Pepsin Trypsin

Goat whey Milk protein

Trypsin, chymotrypsin

Whey protein concentrate (WPC) and heat treated WPC

MBR BMR using polyacrilamide membranes MBR with ultraﬁltration unit

Pepsin, trypsin, chymotrypsin, pancreatin, elastase, carboxypeptidase Protex 6 L from bacillus licheniformis Hydrolytic enzymes

a-lactalbumin and b-lactoglobulin

MBR with two step ultraﬁltration system (30 and 1 kDa)

Production of peptides Recovery of antithrombotic peptides Production of a-lactorphin Production of phosphopeptides Production of polypeptides and rich fraction of small peptides Production of ACE-inhibitory peptides

Whey protein

MBR

Whey protein hydrolisates

MBR with ultraﬁltration unit

Production of whey-protein hydrolyzates Production of emulsifying peptides

[108] [109] [110]

[111]

[112] [113]

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Table 9.9

References

During protein hydrolysis by a membrane bioreactor it has to be considered that an excessive hydrolysis should be avoided because a high content of free aminoacids involves negative effects like bad sensory properties and high osmolarity [114]. This means that to develop the system on an industrial scale, the hydrolytic reaction has to be strictly controlled. Different works were focused on the optimization of process parameters for a continuous production of whey-protein hydrolysates. Guadix et al. [114] developed a MBR with an ultraﬁltration unit (polyethersulfone) where no effects on enzyme activity, due to mechanical shear stress, adsorption to the membrane or enzyme leakage were observed. The effect of temperature on the performance of a batch reactor with an ultraﬁltration unit made of polysulfone material of 8 kDa was analyzed in the hydrolysis of whey-protein hydrolyzates [112]. The experimental data perfectly ﬁt a mechanistic model also proposed in the same article.

9.7 Conclusions

In this chapter the main application of membrane bioreactor and biocatalytic membrane reactor in food with emphasis on the production of functional food is reported. The main aspects were outlined to understand the recent development of the technology and its potential future applications in the ﬁeld. Research efforts are needed to improve aspects such as reproducibility on the large scale, enzyme life-time and immobilized enzyme stability during membrane-cleaning procedures. Technological strategies able to control these parameters are expected to fuel the further development of the biocatalytic membrane reactor on a large scale.

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Ochoa, F. (1998) Enzyme Microb. Technol., 22, 558–567. Bakken, A.P., Hill, C.G., and Amundson, C.H. (1989) Biotech. Bioeng., 33, 1249–1257. Czermak, P., Ebrahimi, M., Grau, K., Netz, S., Sawatzi, G., and Pfromm, P.H. (2004) J. Membrane Sci., 232, 85–91. Petzelbauer, I., Nidetzky, B., Haltrich, D., and Kulbe, K.D. (1999) Biotechnol. Bioeng., 64 (3), 322–332. Neuhaus, W., Novalin, S., Klimacek, M., Splechtna, B., Petzelbauer, I., Szivak, A., and Kulbe, K.D. (2006) Appl. Biochem. Biotechnol., 134, 1–14. Jurado, E., Camacho, F., Luzon, G., and Vicaria, J.M. Enzyme Microb. Tech. (2006) 39, 1008–1015. Gonzalez, R., Ebrahimi, M., and Czermak, P. (2009) The Open Food Science J., 3, 1–9. Tango, M.S.A. and Ghaly, A.E. (2002) Appl. Microbiol. Biotechnol., 58, 712–720. Roy, D., Goulet, J., and LeDuy, A. (1986) Appl. Microbiol. Biotechnol., 24, 206–213. Li, Y. and Shahbazi, A. (2006) Appl. Biochem. Biotechnol., 129–132, 985–996. Gomes, A.M.P. and Malcata, F.X. (1999) Trends Food Sci. Technol., 10, 139–157. Kwon, S., Yoo, I.K., Lee, W.G., Chang, H.N., and Chang, Y.K. (2001) Biotechnol. Bioeng., 73, 25–34. Choundhry, B. and Swaminathan, T. (2006) Appl. Biochem. Biotechnol., 128, 171–183. Shahbazi, A., Mims, M.R., Li, Y., Mishra, M.C., and Karanth, N.G. (2001) J. Chem. Technol. Biotechnol., 76, 529–540. Shahbazi, A., Mims, M.R., Li, Y., Shirley, V., Ibrahim, S.A., and Morris, A.

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(2001) Appl. Biochem. Biotechnol., 73, 529–540. Olmos-Dichara, A., Ampe, F., Ulibearrea, J.-L., and Parelleiux, A. (1997) Biotechnol. Lett., 19, 709–714. Korhonen, H. and Pihlanto, A. (2006) Int. Diary J., 16, 945–960. Fitz-Gerald, R.J., Murray, B.A., and Walsh, D.J. (2004) J. Nutr., 134, 980S–988. Gobbetti, M., Minervini, F., and Rizzello, C.G. (2004) Int. J. Dairy Technol., 57, 172–188. Cheison, S.C., Wang, Z., and Xu, S. (2006) J. Membrane Sci., 286, 322–332. Nakamura, Y., Yamamoto, N., Sakai, K., Okubo, A., Yamamzaki, S., and Takano, T. (1995) J. Dairy Sci., 78, 777–783. Gobbetti, M. (2000) Appl. Environ. Microbiol., 66, 3898–3904. Mannheim, A. and Cheryan, M. (1990) J. Food Sci., 55, 381–385. Bouhallab, S. and Touze, C. (1995) Lait, 75, 251–258. Bordenave, S., Sannier, F., Ricart, G., and Piot, J.M. (1999) Prep. Biochem. Biotechnol., 29, 189–202. Righetti, P.G., Nembri, F., Bossi, A., and Mortarino, M. (1997) Biotechnol. Prog., 13, 258–264. Turgeon, S.L. and Gauthier, S.F. (1990) J. Food Sci., 55, 106–110. Pihlanto-Lapp€al€a, A., Koskinen, P., Piilola, K., Tupasela, T., and Korhonen, H. (2000) J. Diary Res., 67, 9–11. Prieto, C.A., Guadix, E.M., and Guadix, A. (2008) Biochem. Eng. J., 42, 217–223. Gauthier, S.F. and Pouliot, M. (1996) IDF Bull., 311, 31–32. Guadix, A., Camacho, F., and Guadix, E.M. (2006) J. Food Eng., 72, 398–405. Iwasaki, K., Inoue, M., and Matsubara, Y. (1998) Biosci. Biotech. Biochem., 62, 262–272.

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10 Membranes for Food Packaging Alberto Figoli, Erika Mascheroni, Sara Limbo, and Enrico Drioli 10.1 Introduction

The development of new materials for food packaging is a challenge that involves scientiﬁc and technological competences. Consumer needs and socioeconomic problems are the most important driving forces of this process that has, as ultimate goal, the delivery of high-quality and safe food products to the consumer in an efﬁcient manner [1]. All the materials used as food contact materials (FCM) must have speciﬁc and distinctive characteristics. The preliminary requirement is the safety of material. This means that the possible migration of undesirable packaging constituents into the food has to be well known and controlled. The matters of inertness of FCM and packaging reliability are in the domain of law in all the developed countries, where nowadays exist very huge detailed and generally severe regulations on this topic. Other complementary and essential performances concern all those physical and chemical properties that give speciﬁc behavior to the material under conditions of use. For example, physical properties such as gas permeation through package walls, mechanical resistance to environmental stress, sealability, and so on, are very important and useful both for controlling the package-fabrication process and to design the food package able to maintain and guarantee the quality and safety of the product during its shelf life. In fact, environmental factors such as humidity, oxygen, light, and so on (which can induce degradation reactions during storage) should be strictly controlled and in some cases modulated by the packaging material. Synthetic polymers are the materials of choice for many food-packaging applications. They have molecular weights typically between 50 000 and 200 000, an optimum range suitable for shaping the polymers into bags, containers, or other forms that give the adequate protection to food during distribution and storage. The typical properties of common plastic packaging materials are reported in Table 10.1. It is important to highlight that the required packaging protection depends on the product characteristics but not always does the proper protection mean the complete isolation of food from the environment-degradation factors. For example, some fatty foods with long shelf life are sensitive to oxygen and light and, as a consequence, the

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Table 10.1 Main properties of plastic packaging materials. (After modiﬁcation from [1]).

Material PE LDPE LLDPE HDPE PP PS General Impact PVC Unplastic. Plasticized PET PVDC EVOH EVA Nylon

Mechanical property

Moisture barrier

Gas barrier

Use T ( C)

Tough, ﬂexible Tough, extensible Tough, ﬂexible Moderately stiff, strong

High High Very high High

Very low Very low Very low low

50 to 80 30 to 100 40 to 120 40 to 120

Stiff, strong, brittle Tough, strong

Low Low

Low Low

20 to 90 20 to 90

Stiff, strong Soft, extensible Stiff, strong Stiff, strong Stiff, strong Tough, extensible Strong, tough

High Moderate Moderate Very high Low Moderate High

Moderate Moderate Moderate Very high Very high Low High

2 to 80 2 to 80 60 to 200 20 to 130 20 to 150 75 to 65 2 to 120

ideal preservation requires the absence of oxygen inside the package, so a high barrier material to reduce the oxygen entrance and, possibly, no light transmission through the package has to be used. On the contrary, for minimally processed vegetables, the natural interplay between the respiration of the product and the transfer of gases through the packaging can lead to an appropriate atmosphere within package that contributes to maintaining the product freshness during commercialization. In this speciﬁc case, the protection by the packaging is granted by ﬁlms with proper gas permeability that allow the right exchanges between the internal and external sides of the package and not by high barrier materials. In recent years, besides the traditional basic functions of packaging (i.e. protection, communication, convenience, and containment) extra enhanced functions have been sought by the food-packaging sector to meet the consumer demands for minimally processed foods with fewer preservatives, increased regulatory requirements, market globalization and concern for food safety. Active packaging is the main area in which most of recent innovative ideas have been applied to satisfy these needs, broadening and redeﬁning the function of food packaging [1]. Active packaging has been deﬁned as a system in which the product, the package, and the environment interact in a positive way to extend shelf life or to achieve some characteristics that cannot be obtained otherwise. In other words, active packaging is a new generation of packaging materials that can release active compounds (antimicrobial, antioxidants, enzymes, ﬂavors, nutraceuticals, etc.) or absorb undesirable substances (oxygen, ethylene, moisture, etc.) at controlled rates suitable for enhancing the quality and safety of a wide range of foods during extended storage. In this context, food packaging and membrane developers started collaborations in order to evaluate how membrane science could be applied to food packaging area. In

10.2 Application of Membranes in Controlling Gas Permeability

Traditional Food Packaging

- Modified Atmosphere Packaging - Active/smart Packaging

Membranes

Figure 10.1 Fields of applicability of membranes with respect to traditional food packaging.

fact, the wide range of properties required to the packaging gives the idea to design and synthesize the membranes as devices that should contribute to maintaining food quality. An example is the fact that recently in the International Membrane Conferences held in Korea (ICOM06, Seoul), in USA (ICOM08, Honolulu) and France (ICOM09, Montpellier) a speciﬁc session was devoted to food packaging. Moreover, the USA market for nonseparating membranes used in drug delivery, guided tissue regeneration, batteries, food packaging and high-performance textiles was calculated to be $2.8 billion in 2005, more than half the value of the combined market for all the membranes used in separation and nonseparating applications [2]. The deﬁnition of a membrane is not univocal and many attempts have been made to describe it. The most general one may be the following, as reported by Paul and Yampolskii [3]: A membrane is a phase or a group of phase that lies between two different phases, which is physically and/or chemically distinctive from both of them and which, due to its properties and the force ﬁeld applied is able to control the mass transport between these phases. Membranes, both organic and inorganic, are generally classiﬁed based on their morphology as porous, nonporous (dense/tight) and liquid membranes [4]. Depending on the speciﬁc membrane properties (porosity, hydrophobicity/hydrophilicity, pore size, etc.), they can be used as packaging materials in modulating the gasexchange rate between the inside and outside of the package environment (modiﬁedatmosphere packaging) or in actively controlling the release or absorption of speciﬁc compounds to or from the packaged food (active packaging). In this chapter, the use of membranes in food packaging will be analyzed under these two main perspectives, giving results of some examples and potential applications. Figure 10.1 shows the ﬁelds of applicability of membranes with respect to traditional food packaging.

10.2 Application of Membranes in Controlling Gas Permeability

The transport of gas or vapor through a ﬂexible food package (usually made of polymeric ﬁlms) can greatly inﬂuence the quality of the food and, in many situations, the role of packaging that is in the direction of reducing, as low as possible, the gas exchange between the internal or external side of the package. For example, oxygen

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permeability through the package can cause oxidation in lipid foods (dehydrated and processed meat, egg, cheese, fatty foods) that leads to off-ﬂavor production and loss of ﬂavor, color, nutrient value. Water-vapor permeation inside the package can cause moisture gain leading to sogginess or microbial growth in food, while water vapor escaping from the package can cause moisture loss leading to undesirable textual changes in food [1]. The modiﬁed atmospheres (MAP) and the under-vacuum packaging for the storage of nonrespiring products, in fact, require the use of high barrier materials able to reduce the loss of gas during storage, to maintain the optimal atmosphere initially ﬂushed. On the contrary, there are occasions when the transport of gases and vapors is desirable. In modiﬁed-atmosphere packaging of fresh produce (i.e. respiring products), the exchange of oxygen, carbon dioxide and water vapor through the package is necessary to accommodate the respiration and transpiration of the respiring product and to maintain an optimum gas composition in the package. As shown in Table 10.1 polymeric plastic materials cover a wide range of permeability performances but, often, it is necessary to combine the characteristics of one or more materials to reach the desired value. In the ﬁeld of membranes, gas separation can be considered as a major industrial application of membrane technology thanks to the research achievement in the 1960s and 1970s. This progress refers to the development of membrane structures, which allowed to have high ﬂuxes and large surface area modules [5]. The production of asymmetric membranes for reverse osmosis applications by the Loeb–Sourirajan phase-separation process was fundamental for the gas-separation technology growth. Typical asymmetric membranes are made starting from a glassy polymer and the thicknesses of the selective layer are usually between 0.1 mm and 0.5 mm. The main limit of the use of asymmetric dense membranes is that even small defects (pinholes) in the selective membrane, produced during membrane preparation and module manufacture, lead to a decrease of the selectivity of the gas-separation membranes. This problem was partially overcome by coating the membranes with a highly permeable polymer (i.e. silicone rubber), so that the selectivity and ﬂux of the membranes were not signiﬁcantly affected [5]. A different type of membrane used in gas separation is the composite membrane. It is made of a thin selective layer coated on a porous support layer. However, it is difﬁcult to obtain composite membranes with very thin glassy selective layers as those obtained by phase separation. Therefore, composite membranes are usually employed to prepare membranes with a rubber selective layer and the porous substrate to give a mechanical strength. The gas permeation is a chemical-physical phenomenon that concerns, in food packaging, only permeable packages and in particular those consisting of polymeric parts such as plastic ﬁlms, rigid plastics containers, plastic-coated papers and metallized plastic ﬁlms. Usually, membrane processes that utilize polymers include gas separation (e.g., nitrogen generation from air), vapor permeation (e.g., recovery of volatile organic compounds from gas streams), pervaporation (e.g., dehydration of ethanol) and reverse osmosis (e.g., desalination of water) [6].

10.2 Application of Membranes in Controlling Gas Permeability

Permeation is deﬁned as the movement of gases, vapors or liquids (also called permeant substances or penetrant molecules) across a homogeneous material driven by a concentration gradient in the direction from high to low concentration [1]. The permeation of small molecules through a dense polymeric material is described by a solution diffusion model. In fact, gas molecules on the high-pressure side of the membrane dissolve in the polymer, diffuse down the concentration gradient, and desorb on the low-pressure side of the membrane [7]. The permeability coefﬁcient, P is the product of a solubility coefﬁcient, S, and a diffusion coefﬁcient or diffusivity, D: P ¼SD

ð10:1Þ

The permeability coefﬁcient is derived using Ficks ﬁrst law to model diffusion and Henrys law to model adsorption and desorption [8]. Generally, the measure of the ability of a membrane to separate two gases, A and B, is given by the ratio of their membrane permeability or by the product of the diffusion selectivity and the solubility selectivity and it is called membrane permselectivity (a): aA=B ¼

PA ¼ PB

SA SB

DA DB

ð10:2Þ

Most of the recent research in membrane science has focused on developing membrane materials with a better balance of selectivity and permeability as this seems the most likely route for expanding the use of this technology also in food packaging. Figure 10.2 shows the typical trade-off between the CO2/N2 and O2/N2 selectivity and CO2 and O2 permeability, respectively, for a vast number of membrane materials [9].

Figure 10.2 Upper-bound correlation for CO2/N2 and O2/N2 separation. (reprinted from [9], with permission of Elsevier.)

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The graph suggests that exists an inherent relationship between the selectivity of a polymeric material and its permeability. The lines above which no data points exist are called the upper bounds. Over the past 20 years these upper bounds shifted to higher values, however, the direction still remains valid: a high permeable polymer material frequently has a low selectivity and vice versa. Therefore, the overcome of such upper bound motivates material scientists to develop new concepts to realize high productivity as well as high-selectivity membranes. The relationship between the selectivity and polymeric material is of fundamental importance also in designing membranes for speciﬁc food packaging application. 10.2.1 Membranes in Modified-Atmosphere Packaging

The shelf life of horticultural products and minimally processed fruits and vegetables (i.e. fruits and vegetables that have attributes of convenience and fresh-like quality) is limited by respiration, transpiration and enzymatic activity of the living tissue especially after harvest processing and, at the same time, to proliferation of spoilage and pathogenic micro-organisms. To reduce the effects of these factors it is possible to act on processing or, more usually, on packaging. Modiﬁed-atmosphere packaging (MAP) is effective in prolonging shelf life by decreasing O2 and increasing CO2 concentrations in the package atmosphere that successively changes as a consequence of respiratory O2 uptake and CO2 evolution of packaged product (respiration rate) and gas transfer from the packaged ﬁlms. In other words, the basic principle of MAP is that a modiﬁed atmosphere can be created passively by using properly permeable packaging materials, or actively by using a speciﬁed gas mixture together with permeable packaging materials. A proper combination of product characteristics, ﬁlm permeability and ﬁlm selectivity results in the evolution of an appropriate atmosphere within packages [10, 11]. The aim is to create an optimal gas balance inside the package, where the respiration activity of a product is as low as possible and on the other hand, the oxygen concentration and carbon dioxide levels are not detrimental to the product. In fact, it has been demonstrated that the respiration rate depends on the concentration of O2 and CO2 in which the product is stored after harvesting. Based on these results, several mathematical models have been developed [12, 13]. In particular, the one based on the principles of enzyme kinetics (the Michaelis–Menten kinetic equation) which also take into account the inhibition due to CO2 is reported here [13]. The suggested form for the rate of O2 uptake is the following: rO 2 ¼

WRm pO2 pCO2 pO 2 km þ 1 þ Ki

where, pO2 and pCO2 are the partial pressures of O2 and CO2, W is the mass of the product, and Rm, km and Ki are the temperature-dependent parameters speciﬁc for a particular type of product. A similar equation can be reported also for the rate of CO2 production, rCO2 .

10.2 Application of Membranes in Controlling Gas Permeability

The ratio of CO2 produced to O2 consumed is known as the respiratory quotient, that is, RQ ¼

rCO2 rO2

This ratio can range from 0.7 to 1.4 depending on the substrate and metabolic state [14]. Actually, the most difﬁcult task in manufacturing raw ready-to-use or ready-to-eat fruit and vegetable products of good quality and possessing a shelf life of several days is to maintain such optimum concentration of O2 and CO2. The main problem is that only a few packaging materials on the market are permeable enough to match with the respiration of fruits and vegetables. Moreover, with fresh respiring products, it would be advantageous for the product shelf life retention to have ﬁlm permeability increased by temperature, at least as much as the respiration rate increases in order to avoid anaerobic conditions. Unfortunately, the permeation rates of most packaging ﬁlms are only modestly affected by temperature [15]. One approach of extending the shelf life of fruits and vegetables by membranes is reported by Paul and Clarke [16]. They worked on modeling the performance of packages containing respiring products that have both a permselective membrane (asymmetric dense membrane, rubber or glassy type) and perforations (nonselective membrane) as shown in Figure 10.3. Such packaging has the function to regulate the permeation of oxygen and carbon dioxide, respectively into and out of the food packaging, to reach a steady state between the respiring product and the external atmosphere of the packaging. In fact, usually selective membranes are able to create the gas compositions needed only for a certain group of products, since they are too selective in permeation of CO2 relative to O2, while perforations are not selective. The model calculations showed that a wide range of gas optimal atmospheres for many fresh products, such as broccoli, mangoes, cauliﬂower, and so on, can be created using a semipermeable membrane together with the perforations that provide a nonselective permeation of gases between the air outside the package and the gas mixture inside the package. In another article, the same authors [17] extended the same concept from small disposable retail packages to reusable large-scale containers for storage and shipping

CO2

Non Selective Perforation

O2 rCO2

Selective Membrane Dense layer Porous Substrate

product

rO2

Plastic Film

Figure 10.3 Schematic drawing plastic food package with a selective membrane patch and nonselective perforations or holes. (After modification from Ref. [16], with permission of Elsevier.)

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Controlled air flow Product storage chamber

Air flow in large excess

Reject flow Non-selective membrane

rCO r 2 O

2

product Selective membrane

Figure 10.4 Scheme of the modified-atmosphere packaging concept of a large-scale, reusable container for storage or shipping of food products combining the use of a selective and nonselective membrane. (After modification from Ref. [17], with permission of Elsevier.)

of food products. The concept is shown in Figure 10.4. Also in this case a selective membrane is used in combination with a nonselective membrane acting in parallel. The relative amount of gas exchange (CO2 and O2) through the nonselective membrane can be adjusted by varying the volumetric air feed rate to its upstream surface that will, in turn, correct the steady-state composition in the product chamber. Therefore, a desired atmosphere can be created by regulating the air feed rate that can reduce the selectivity of the CO2 to O2 generated by the selective membrane to enable operation at higher CO2 concentrations. Furthermore, the mathematical model developed shows that the lower CO2 concentration in the product storage chamber is largely dependent on the selective membrane, whereas, the higher CO2 concentration is mostly dependent on the choice of the nonselective membrane. The region between these lower and upper CO2 concentration limits required by a wide range of products can be reached by adjusting the air feed rate over a range of 1–2 orders of magnitude. In another study, Torchia et al. [18] reported the application of a novel polymer material, the modiﬁed polyaryletheretherketone (PEEKWC), to food packaging. This polymer has excellent chemical, thermal and mechanical properties, and it has the advantage, compared to traditional PEEK, to be soluble in several common organic solvents, facilitating the asymmetric (dense and porous) membrane preparation by phase inversion [19–22]. Different types of membranes with pure PEEKWC and PEEKWC modiﬁed with poly-a-pinene (PaP), PEEKWC/PaP, were prepared by solvent evaporation. The ﬁlms produced, PEEKWC and PEEKWC/PaP, were characterized in terms of gas and vapor transport (PO2 and WVTR) and the results obtained were comparable to those of commercial food-packaging polymers. The potentiality of these materials was tested for MAP of fresh products, such as fruits and vegetables. The selectivity was calculated as the ratio between the permeability values of the O2 and CO2 species that obey to the Arrhenius law in the analyzed temperature range. The activation energy values for the permeation of O2 and CO2, respectively, were also calculated. The separation performance, CO2 and O2 permeability, of the polymer as well as the CO2/O2 selectivity changed at higher additive concentration as reported in

10.2 Application of Membranes in Controlling Gas Permeability Table 10.2 Oxygen and carbon dioxide permeability and activation energy in polymeric films at 5 C. (After modification from [23]). (Reprinted from [18], with permission of Chiriotti Ed.).

Polymers

PEEKWC PEEKWC/PaP (80/20) PEEKWC/PaP (50/50) Silicon rubber Natural rubber Polybutadiene Poly(butadiene-styrene) LDPE HDPE PA 6 Saran PET Cellulose acetate PVC

PO2 (Barrer)

PCO2 (Barrer)

E O2 P (kJ mol1)

E CO2 P (kJ mol1)

a (CO2/O2) (5 C)

0.36 0.30 0.12 100 8.7 7.7 6.7 1 0.014 0.007 0.0004 0.013 0.34 2.1

2.4 1.7 0.56 65 61.4 71 59.6 7.4 0.07 0.035 0.0042 0.045 2.9 12.7

21.7 21.7 37.4 8.4 31.4 29.7 30.5 43.1 35.1 43.5 66.5 26.8 20.9 36.9

11.3 15.3 31.3 — 25.5 21.8 23.8 34.2 30.1 40.6 51.5 25.9 29.7 27.6

6.7 5.7 4.7 6.5 7.1 9.2 9.2 6.7 4.8 4.7 10.2 3.4 8 6.1

Barrer ¼ 1010 (cm3 (STP) cm/cm2 s cmHg).

Table 10.2. The CO2/O2 selectivity, measured through the PEEKWC ﬁlms, is comparable to that of LDPE and silicone rubber ﬁlms. The O2 permeability is about one third higher than LDPE but two orders of magnitude lower than silicone rubber. PEEKWC/PaP 50/50 ﬁlms were characterized by a CO2/O2 selectivity of 4.7 similar to the one of HDPE and Nylon 6. In this case, the O2 permeability of PEEKWC/PaP 50/50 results 8.5 and 170 times, respectively, higher than that measured in the previous commercial polymers. Oxygen and carbon dioxide permeability of PEEKWC were similar to cellulose acetate ones In Table 10.3, an increase of the activation energy for CO2 permeation is also observed when the additive amount rises. On the other hand, the activation energy, calculated for oxygen, is constant up to 20% PaP concentration, it increases signiﬁcantly at 30% of additive, remaining almost constant up to 50/50 PEEKWC/PaP ratio. The CO2/O2 selectivity, calculated for pure and added PEEKWC ﬁlms, suggests its use in food packaging only for a restrict number of fresh products packaging, see Table 10.3 [18]. Concerning the permeability of PEEKWC ﬁlms, with and without additive, it is still too low with respect to the majority of fresh fruit and vegetables studied. However, this limitation can be extended generally to all polymeric materials. Only a few packaging materials on the market are permeable enough to match with the respiration of fruit and vegetables. Most ﬁlms do not result in optimal O2 and CO2 atmospheres, especially when the product has a high respiration rate. Therefore, the development of novel membranes, that is, loaded with nanostructure material, or, alternatively the presence of nonselective membranes, as previously discussed [16], carefully distributed in the packaging, could extend the use of PEEKWC or other polymeric material to a broader number of products.

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Oxygen and carbon dioxide permeability (PRO2 and PRCO2) and selectivity (aR) required for various fruits and vegetables in typical market size packages [23] compared with PEEKWC based films. Film thickness 25 mm and 4 C. (Reprinted from [18], with permission of Chiriotti Ed.).

Table 10.3

Fresh products

PRO2

PRCO2

aR (CO2/O2)

Strawberry Bruxelles sprouts

22.4 12.7

23.9 76.6

1.1 6

Mushrooms Lettuce Turnip

10.6 4.5 2.2

13.1 42.94 7.6

1.3 9.5 3.5

Carrot

1.54

5.7

3.7

Apple

1.51

9.5

6.3

Celery Green Pepper

1.26 0.67

4.04 4.03

3.2 6

Tomato Blackberry

3.65 6.4

19.2 20.1

a (CO2/O2) of PEEKWC films

(90/10) and (80/20) PEEKWC/PaP (a ¼ 5.7), PEEKWC (a ¼ 6.7)

(70/30) PEEKWC/PaP (a ¼ 4) (70/30) PEEKWC/PaP (a ¼ 4) (90/10) and (80/20) PEEKWC/PaP (a ¼ 5.7), PEEKWC (a ¼ 6.7) (90/10) and (80/20) PEEKWC/PaP (a ¼ 5.7), PEEKWC (a ¼ 6.7)

5.3 3.14

Barrer ¼ 1010 (cm3 (STP) cm/cm2 s cmHg).

10.3 Membranes as Devices for Active Food Packaging

The concept of controlled delivery using membranes has been mainly applied in the medical ﬁeld in which a membrane is used to moderate the rate of delivery of drug to the body. The application of the membrane differs depending on the type of device employed [24], as shown in Figure 10.5: (a) the membrane has the function to control the permeation of the drug from a reservoir to achieve the desired drug-delivery rate (reservoir system); (b) the drug is dispersed or impregnated into the membrane material and it slowly dissolves or degrades in the body. In this case, the drug

Figure 10.5 Illustration of the membrane controlled-release devices: (a) reservoir system; (b) matrix system, (c) osmotic system.

10.3 Membranes as Devices for Active Food Packaging

delivery is controlled by a combination of diffusion and biodegradation (matrix system); (c) the drug is released using the osmotic pressure developed by diffusion of water across a semipermeable membrane into a salt solution that pushes it out (osmotic system). These concepts have also been extended to other areas of interest to control the delivery of agrochemicals (pesticides), household products (fragrances) and active agents for food-packaging applications (antimicrobials, antioxidants, aroma compounds). Active packaging has been deﬁned as a system in which the product, the package, and the environment interact in a positive way to extend shelf life or to achieve some characteristics [25]. A new challenge in the food industry is the current trend in consumer demands for minimally processed, easily prepared and ready-to-eat fresh products. Traditional preservation of such products, in which the preservative is added directly to the food, has limited beneﬁts. In fact, the active substances are neutralized on contact or diffuse rapidly from the surface, where contamination primarily occurs, into the food mass [26]. Moreover, the addition of large amounts of antimicrobials directly to ready-to-eat products can inﬂuence the taste, while low amounts result in a short shelf life. The controlled release of agents obtained by an active food-packaging system can be generally considered the solution to preserve the quality and increase the storage time for ready-to-eat perishable foods [27]. All the active packaging technologies involve some physical, chemical, or biological action for generating interactions between the package, the product, and the package headspace to increase the shelf life of foods. In addition, they can be divided into categories of absorber, releasing system and other systems [26]. The actual techniques can be summarized as follows: . . . . . .

addition of sachets/pads containing volatile agents; incorporation of volatile and nonvolatile agents directly into the polymers; coating or absorbing agents onto polymer surfaces; immobilization of agents to polymers by ion- or covalent linkages; use of polymers that are inherently antimicrobial; multilayer ﬁlms with active layer.

In particular, multilayer ﬁlms are an interesting solution for active packaging and membranes could have a key role in these multilayer structures. In monolayer dense ﬁlm in fact only a part of the preservative is released [28, 29] and higher concentrations of antimicrobial agents than usually needed have to be loaded in these ﬁlms to preserve the food. Moreover, the release rate of the active compounds is not easily controlled. The multilayer ﬁlms presented in the literature are usually produced by coextrusion of dense ﬁlm or coating of an active thin layer on the polymer surface; the active layer functions both as reservoir and as release control of the active substance [30–32]. In the literature, only few works report the preparation of multilayer ﬁlms, having an outer barrier dense layer, an active agent-containing matrix layer and a release-control layer. The control layer is the key layer to control the initial time-lag period and the ﬂux of penetration of the active agents. In these multilayer structures the membranes

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are suitable devices. Han et al. [33], for example, suggested the use of such multilayer structure for antimicrobial-release packaging systems. Another study on multilayer ﬁlms concerns the controlled release of a volatile antimicrobial compound, the allylisothiocyanate (A.I.T.C). The multilayer ﬁlm is made of (a) a tie-layer, cyclodextrins containing the A.I.T.C and (b) perforated membrane, within a ﬁne powder of silica gel, which is in contact with the food product [34]. Figoli et al. introduced the use of the asymmetric porous membrane in controlledrelease food packaging, produced by nonsolvent-induced phase separation (NIPS) technique [35, 36]. They reported the development of an antimicrobial food packaging ﬁlm based on the use of membranes, with modulate porosity, as a controlling release system. The multilayer ﬁlm was made of three layers: an outer dense layer to control the exchange rate of gases between the external and internal environment of the food packaging, an intermediate adhesive tie-layer which has also the function of reservoir of antimicrobials, and the porous membrane layer, made by phase inversion, that controls the release of antimicrobials to the food. In particular, its properties (porosity and morphology) can be properly tailored by changing the phase-inversion process conditions. In this case, the investigated polymer was the modiﬁed polyaryletheretherketone (PEEK-WC), already widely used in membrane preparation [20, 22]. However, the proposed process can be extended also to other polymer traditionally employed in food packaging. The multilayer ﬁlms were prepared as shown in Figure 10.6. All separate layers of the multilayer ﬁlm could be cast subsequently on one another without removing the dense ﬁlm from the glass substrate. The method to produce the three layers is presented, as follows: 1)

The dense PEEK-WC/poly-a-pinene (p-a-p) (different ratio, from 100/0 to 50/50) layer, used as substrate for the multilayer ﬁlm, was prepared by casting solution. A dense PEEK-WC/p-a-p ﬁlm was formed on the clean glass substrate after the solvent was evaporated. The addition of p-a-p has the double function to increase the afﬁnity of the dense ﬁlm with the second layer and to modify the transport properties of the ﬁlm itself. 2) The second layer was made of p-a-p with and without oxalic acid (0.5, 10 and 25 wt%). The starting solution was stirred and cast at 0% RH and 70 C in a climate chamber. Immediately after casting, the formed double layer ﬁlm was removed from the climate chamber and allowed to cool at room temperature.

Figure 10.6 Schematic representation of the multilayer film casting process developed.

10.3 Membranes as Devices for Active Food Packaging

3)

A porous PEEK-WC ﬁlm was then cast on the double-layer ﬁlm previously prepared. The porous membrane layer was prepared by dry-wet phase inversion. A casting solution was produced with different concentrations of PEEKWC in N,N diethylacetamide (DMA) (15, 19 and 23 wt%). The ﬁlms were cast in a climate chamber at 50% RH and 20 C. The porosity and morphology of each membrane were varied changing the time of exposure to air before precipitation (45 s/240 s.) and the water-bath temperature (0 C and 40 C). The multilayer ﬁlm was removed 5 min after immersion from the coagulation bath. The same PEEKWC membranes have also been prepared by casting the solution directly on the glass substrate to evaluate the effect of the polya-pinene substrate, with and without oxalic acid, on the properties (i.e. morphology, porosity) of the membrane. The different membranes obtained by varying the concentration of oxalic acid and polymer, are illustrated in Figure 10.7. In particular, the increase of the polymer concentration (from 19 to 23 wt%) produces an asymmetric dense membrane.

The ﬁnal antimicrobial multilayer ﬁlms have also been examined by scanning electron microscopy, SEM. The individual layers of the ﬁlm, indicated by the arrows, can clearly be distinguished in the picture (Figure 10.8). The release rates of oxalic acid from the multilayer ﬁlm was determined by bringing into contact the porous membrane side with distilled water and monitoring the change of pH with time. The results proved that the release rate depended strongly on the phase inversion processing conditions and on the compounds used in the preparation of porous membrane layer (air exposure time, water-bath temperature, polymer concentration, oxalic acid concentration). In particular, the oxalic acid release increased with decreasing the coagulation bath temperature (from 40 to 0 C) and when the operating temperature was increased from 5 to 25 C [35, 36].

Figure 10.7 SEM cross-sections of the different membranes (release layer) obtained by changing the concentration of oxalic acid (from 5 to 25 wt.%) and PEEKWC polymer (from 19 to 25 wt.%).

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Figure 10.8 Cross-section of the multilayer film (SEM magnification of 400).

Based on these results, a production line for scale production of the multilayer ﬁlm was proposed as shown in Figure 10.9. The polymer is extruded to produce a dense ﬁlm with barrier or speciﬁc gas or water-vapor properties, then, a commercial tie-layer resin, loaded with a speciﬁc antimicrobial, is cast on the dense ﬁlm. Finally, the polymer solution is spread on the adhesive layer and brought into contact with the coagulation bath (i.e. water) that will determine the formation of the asymmetric porous membrane. Another example reported in the literature is that of Altinkaya et al. who presented the incorporation of lysozyme [37] and natural antioxidants [38], such as L-ascorbic acid and L-tyrosine, into cellulose acetate (CA) asymmetric porous structures by phase-inversion technique. In order to achieve controlled release of the active compounds studied, the ﬁlms structure was modiﬁed by changing the morphology (from porous to dense) tailoring the composition of the initial casting solution. In particular, the ﬁlms were produced using the dry phase-inversion technique. The polymer was dissolved in a mixture of acetone and water and, then, cast on a support and exposed to an air stream. Different morphologies were obtained by changing the phase-inversion processing conditions such as evaporation temperature, relative humidity, wet casting thickness as well as the composition of the membrane-forming solution. In the case of lysozyme [37], the highest release rate and antimicrobial activity were obtained with the ﬁlm prepared with 5% CA solution including 1.5%

Figure 10.9 Production line suggested for the fabrication of the multilayer film with the asymmetric membrane as the antimicrobial controlled-release system.

10.3 Membranes as Devices for Active Food Packaging

lysozyme. At higher CA concentration (15%) the porosity of the ﬁlm was reduced with a consequent decrease of the release rate. The diffusion of lysozyme in CA, porous and dense, ﬁlms was 4.17 1010 (cm2 s1) and 1.50 1010 (cm2 s1), respectively. The mechanical properties of the ﬁlms were evaluated also in terms of tensile strength, % elongation at break and Youngs modulus. The tensile strength, Youngs modulus and elongation at break of the ﬁlms increased with increasing CA concentration due to reduced pore sizes and porosity of the ﬁlms. The incorporation of lysozyme into the ﬁlms prepared with 5% and 10% CA solution did not determine any change in the mechanical properties with respect to the ﬁlms without lysozyme. In contrast, the ﬁlm prepared with 15% of CA loaded with lysozyme showed a signiﬁcant reduction in tensile strength and elongation at break values. Also, in the case of the loading of low molecular weight natural antioxidants [38], such as L-ascorbic acid and L-tyrosine, the diffusion rate through the ﬁlms was reduced by increasing the CA concentration in the casting solution. The use of the porous or dense structure in contact with food environment and the different CA concentration of the made ﬁlm are the main factors responsible of the release rate of these active compounds. The diffusion rate of L-ascorbic acid was 3.33 1010 (cm2 s1) and 1.67 1010 (cm2 s1) in porous and dense structures, respectively, while for L-tyrosine was 1.00 1010 (cm2 s1) and 0.8 1010 (cm2 s1) in porous and dense structures, respectively. Also the mechanical properties of the ﬁlms increased signiﬁcantly on increasing the CA concentration, due to the fact that the ﬁlms produced had a lower porosity, pore size and they were practically dense. Figoli et al. [39], recently illustrated the advantages of using microencapsulation as a promising technology for protecting the natural active substances from the stresses and damages that can occur during food-package manufacturing and for improving the active-agent distribution. Thanks to these effects and according to their structure, the microcapsules could better control the release of the active substances and promote the interaction of the ﬁlm with the active substances carrier. In this work, bio-microcapsules of chitosan have been developed using a system that combines the membrane process concept with the phase-inversion technique using a monoporous polymeric ﬁlm [40]. This technique permitted the formation of monodispersed biopolymer droplets that were then cross-linked with a natural additive adapted for this polymer structure, and that enhanced the water resistance of chitosan itself. The capsule size and morphology were adjusted by changing the ingredient parameters such as the cross-linking concentration and tailored with the pore diameter of the monoporous ﬁlm employed. Furthermore, two different types of natural antimicrobial were included in the capsules enabling loading both during their production and after the droplet formation. The chemicalphysical analysis of the new chitosan microcapsules was carried out by means of optical microscopy, SEM and EDX. The chitosan capsules produced are shown in Figure 10.10. The antimicrobial activity of the microcapsules was assayed by turbidimetric methods against Staphylococcus aureus selected as a pathogen micro-organism, which

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Figure 10.10 Optical microscope image of chitosan microcapsules without (a) and with (b) antimicrobical compound.

may be present in fresh food. The results showed that the addition of the antimicrobials enhanced the antimicrobial effect of chitosan itself and the growth of Staphylococcus Aureus was totally inhibited.

10.4 Conclusion

In a period in which consumers are demanding higher-quality foods and changes in retailing practices (such as market globalization resulting in longer distribution of food), or the consumers way of life (resulting in less time spent shopping for fresh food at the market and cooking), the major challenge to the food-packaging industry is the development of new packaging concepts that extend shelf life while maintaining and monitoring food safety and quality. In this context, membranes can play an important role and, even if their potentialities have not been completely exploited, some promising membrane features in food packaging have been reported and discussed in this chapter. Their use has been addressed toward two main perspectives: (a) modulating the gas exchange rate between the inside and outside of the package environment (such as some applications of modiﬁed-atmosphere packaging) and (b) actively controlling the release or absorption of speciﬁc compounds from the packaged food (active packaging). In particular, the possibility of tailoring the membrane morphology, porosity and properties allows an extension of their use to a broad range of food-packaging purposes. Furthermore, the growing environmental awareness coupled with the inexorable rise of pre-packaged disposable meals has directed the research to the development of environmentally friendly packaging materials with biodegradable properties, preferably with components from natural sources rather than from petrochemical materials. Even if up to now biopolymers have been slow to reach commercial maturity, due to their high costs and less optimal physical properties than conventional plastics, things are changing, and new large-scale production systems are bringing down the costs of biodegradable polymers, and sophisticated polymerization and blending techniques are improving the material properties. In this continuously evolving scenario, the design and production of innovative packaging is of fundamental importance and membranes can have a key role and actively contribute to this innovation.

References

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Index a

c

accumulation 30f. adsorption 36, 126, 137, 209, 227 aggregation 50, 155 antioxidants 167, 205, 237 – activity 205 – total antioxidant activity (TAA) 180f.

cake formation 32, 36 carbonation 22 casein micelles – concentration 25, 28f., 35, 45 – rejection 57, 59 – separation 47, 55 – skim milk 28f. – transmission 48 cell immobilization 211 centrifugation 25, 27, 50, 66, 81 – wine production 106 CFD (computational ﬂuid dynamics) 35f., 132, 134f., 138 cheese – curding 26 – partial concentration 8 – preconcentration 8 – preservation 4 – processing 8f., 51 – standardized cheese milk 9 – total concentration 9 – whey 30 chromatography – fractionation 48 – ion-exchange membrane 66 Clapeyrons law 175 clariﬁcation 9 – beer 9ff. – fruit juice 16f. – MF/UF 105ff. – vinegar 15 – wine 13, 209 cleaning – chemical 33 – electrodialysis systems 99 – in situ 82 – methods 31

b backtransport models 32, 35 – backpulsing 28, 32f. – backwashing 33 bacteria removal – milk 3f., 46 – skim milk 27 beer – clariﬁcation 9ff. – dealcoholization 9, 11f. – maturation 11 – membrane bioreactor 212f. – production 10ff. bioactive compounds 217f., 224 biocatalyst 201ff. – esterases 206 – GRAS (generally regarded as safe) 214 – laccase 205 – lipases 206f. biocatalytic membrane reactor, see membrane bioreactor biological oxygen demand (BOD) 5 biologically active compound 167 biotransformation 201 brine 4, 16 – boundary layer 174 – treatment 8 Brownian motion 152 by-product 201f.

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– sequences 1, 110f. COD (carbon oxygen demand) – dairy efﬂuent 50f. – dairy process water 66ff. – dairy wastewater 94 – reduction 52 concentration factor, see wine concentration polarization 36, 52, 60, 79, 175, 186, 189 conductivity – dairy-process water 66f. – ionic concentration solutions 88f. – wine 113 costs – energy 2, 8, 30, 47, 62, 87, 168f. – equipment 8, 63 – membrane replacement 17, 170 – operating 10 – packing 168 – production 8 – shipping 168 – storage 168 – water 9 cross-ﬂow velocity 38 cryoconcentration 168, 194 crystallization – cold 106 – induction time 114f. curding 25f. current – density 79f. – efﬁciency 87f.

d dairy efﬂuents 51f. Dean vortices 33 deionization degree (DEID) 113f. desalination 1, 75, 95f. diaﬁltration (DF) 5, 13 – milk 48 – whey demineralization 8 – whey proteins fractionation 49 – wine 13f. diffusion – Brownian 37 – coefﬁcients 37, 175, 227 – rate 237 – selectivity 227 – shear-induced 35, 37 direct osmosis (DO) 194 distillation – applications 177ff. – ﬂux 174f. – fruit juice concentration 171ff.

– membrane (MD) 168, 170f., 183ff. – membrane osmotic (MOD) 190f. – osmotic (OD) 22, 121f., 170ff. Donnan exclusion mechanism 22, 75 droplet – break-up 137 – coalescence 130, 152 – force-balance model 135, 137 – formation 134ff. – generation unit (DGU) 133, 151f. – lipid-coated ice droplet hydration method 147 – -size 131, 134, 139, 143 – -size distribution (DSD) 129, 148, 151ff. – stabilization 142ff. – torque-balance model 135, 137

e electrodoalysis (ED) 1f., 8, 21f. – acidiﬁcation 125f. – applications 82ff. – cost analysis 79ff. – dairy industry 88ff. – deacidiﬁcation 84f., 168 – demineralization 30, 50, 84, 88ff. – electrochemical coagulation 97f. – electrodeionization 96f. – electroreduction 98f. – food industry 75ff. – juice industry 85ff. – lactic-acid production 90ff. – process 76f., 112 – protein fractionation 94ff. – reversal (EDR) 75 – stack 78, 81f. – sugar industry 85ff. – system design 79ff. – tartaric stabilization 105f., 111ff. – wine 83ff. emulsiﬁcation – coefﬁcient of variation (CV) 132 – cross-ﬂow membrane (XME) 129ff. – dead-end membrane (PME) 130f., 134, 137, 144, 150 – dispersed phase 138f., 141, 143, 146f. – encapsulation active molecules 149ff. – low-shear processing 156f. – membrane (ME) 129, 134, 136, 139ff. – microchannel (MCE) 129, 133f., 138f., 141ff. – rotating membrane (RME) 130ff. – vibrating membrane (VME) 132 emulsiﬁer 7, 45, 206

Index – hydrophile–lipophile balance (HLB) 142, 148 – hydrophilic 149 – hydrophobic 146, 149, 151 – proteins 142 – small-molecule 137 emulsion – ﬂocculated networks 155 – monodisperse 133f., 139, 143, 146, 150f. – oil-in-water (O/W) 130, 132, 140ff. – polydisperse 153, 155 – semiliquid 152 – stability 152ff. – water-in-oil (W/O) 130, 144ff. – water–oil–water (W/O/W) 147ff. – yield stress 154 enological processes 106, 119ff. enzyme – dosage 204 – immobilization 209, 213 enzyme membrane reactors (EMRs) 167f. enzyme – oxidative 209 – stability 209 enzymatic – activity 208f., 213 – cell-free reactors 212 – coagulation 8 – hydrolyse of starch 203 – pulping 168 – synthesis 206 evaporation 5, 17 – compression 91 – ﬂux 172, 175f. – multistage 168 – solvent 143, 230 – thermal 194 – vacuum (VE) 50, 91, 115 extraction – liquid–liquid 23 – mass agent 202

f fermentation – aerobe 15 – beer 9ff. – wine 12ff. ﬁltration – cross-ﬂow 3, 9f., 21, 38, 45f., 48 – dead-end 4, 9, 28, 106 – diatomaceous-earth 105f. – dynamic 45, 51ff. – kieselguhr 4, 11, 106 – shear-enhanced 51f.

– sterile 12, 106 ﬂow – resistance 82 – reversal 33 – -through models 35 – turbulent 2 ﬂux – critical 2, 31f., 35 – decrease 30f. – steady-state 36, 188 foaming 6, 22 food packaging – active packaging 224f., 232ff. – controlled-release polymers 226, 232ff. – ﬁlms 229ff. – food contact materials (FCM) 223 – modiﬁed-atmosphere (MAP) 225f., 228ff. – multilayer ﬁlms 233ff. – synthetic polymers 223ff. – traditional 225 – under-vacuum 226 force – drag 135 – interfacial tension 134f., 137 fouling 1, 26, 30f., 46, 50, 107 – degree 110 – electrodialysis systems 87, 99 – internal 57, 63 – layer 30 fractionation 5, 35f. – milk-fat globules 27 – sugar/organic acids 116 – whey proteins 48f., 52 free biocatalysts membrane bioreactor, see membrane bioreactor fruit juice – clariﬁcation 16, 167 – concentration 167f., 176 – deacidiﬁcation 169 – depectinization 167 – integrated processes 167ff. – production 16ff. – quality 180, 182 – single-strength 168 – stabilization 167 functional food – ﬂavor 201 – ingredients 203 – membrane bioreactors 202ff. – nutraceutical 203 – nutrapharmaceutical 208 – odor 201 – prebiotic 203 – probiotic 203

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g gas-exchange rate 225 gas separation (GS) 168 gel ﬁltration mode 32 GMP (K-casein-glycomacropeptid) 7

h heat – transfer coefﬁcient 187 – treatment of milk 7, 27 hydrogel coating 173

i ice-cream manufacturing 5 interaction – chip surface–emulsiﬁer molecule – electrostatic 139 – hydrodynamic 135, 155 – nonattractive 139 – particles–liquids 37 – protein–ZrO2 49 interface – liquid/membrane 190 – oil/water 208 – vapor/liquid 171, 183 interfacial – area 170 – rheology 137 – tension 134f., 137f. isoelectric point 131

139

l lattice-Boltzmann method 37, 134 Loeb–Sourirajan phase-separation process 226

m macromolecular 106f. mass transfer – boundary layer 186 – control 202 – distribution coefﬁcient 22 – gas–liquid 22 – liquid–liquid 22 membrane – anion-exchange 76f., 91, 93 – area 32, 62 – asymmetric 26, 226, 229f., 234, 236 membrane bioreactors – beer processing 212f. – biocatalytic membrane reactor (BMR) 201ff. – continuous-ﬂow stirred-tank membrane reactor 209

– ethanol production 212f. – fat processing 206ff. – free biocatalysts membrane bioreactor (MBR) 201ff. – fruit-juices production 213 – lactose hydrolysis 215ff. – milk processing 214ff. – oil processing 206ff. – pectin hydrolysis 210f., 213f. – protein hydrolysis 215 – starch processing 203ff. – stirred-tank reactor 207 – sugar processing 203ff. – two-separate phase membrane bioreactor 207 – wine 208ff. membrane – bipolar 76f., 87, 90ff. – cation-exchange 76f., 79, 87, 91, 93 – ceramic 3f., 26, 29f., 47f., 59, 130, 172f. – chemistry 99 – composite 30, 226 membrane controlled-release devices 226, 232f. membrane distillation (MD) – applications 187ff. – air-gap (AGMD) 184 – direct contact (DCMD) 184, 187ff. – operating parameters 186f. – sweeping gas (SGMD) 184 – vacuum (VMD) 184, 187, 188ff. membrane – ﬂux 29, 226 – hydrophilic 1, 130, 145, 169, 172, 208 – hydrophobic 130, 145, 169, 173, 183, 208 – ion-exchange 1, 22, 30, 50, 75ff. – lifetime 82, 170 – liquid 225 – metallic 130 – modiﬁcation 37 membrane modules – circular polymer 66 – helically wound hollow-ﬁber 174 – membrane distillation (MD) 185ff. – microporous polymer hollow ﬁber 174ff. – osmotic distillation (OD) 172ff. – plate-and-frame 10f., 174, 185 – rotating-disk 58ff. – shell-and-tube 186 – spiral-wound (SW) 4, 17, 48, 51, 66, 185, 202 – tubular 10, 49, 59, 185, 202 – VSEP 60f., 63f., 66

Index membrane – multichannel 47 – nonporous 21, 225 – nonselective 229f. – nuclepore 36 – osmotic distillation (OD) 172ff. – performance 31, 34 – permselectivity 227, 229 – polymeric 21, 30, 47f., 75, 185 – porosity 136, 169, 173, 185, 225 membrane processes – applications in food industry 17ff. – continuous 10 – ED, see electrodialysis – hybrid 23 – integrated 167ff. – MF, see microﬁltration – NF, see nanoﬁltration – PV, see pervaporation – pressure-driven 1f., 186 – RO, see reverse osmosis – three-stage 66 – two-stage 49 – UF, see ultraﬁltration membrane – radius 58 – regeneration 110 – SCT 47 – selectivity 32, 65, 121, 202, 226ff. – semipermeable 229 membrane separation – models 17, 34f. – of components 27ff. membrane – SPG (Shirasu-porous-glass) 130f., 140f., 145, 149 – supported-liquid 170, 191ff. – surface area 108f. – surface morphology 31f. – thickness 173, 189 – track-etched 26 – volumetric productivity 32 – zeolite 21 microdispersions 134 microencapsulation 237f. microﬁltration (MF) 1f., 26 – bacteria removal 3f., 46 – beer clariﬁcation 11 – cross-ﬂow 11 – /electrodialysis (ED) 80 – milk 46 – osmotic distillation (OD) 177, 179, 181f. – polysaccharide removal 113f. – powder milk 56

– productivity 107f. – skim milk 58 – spore removal 46 – tangential MF/UF 105f. – wine clariﬁcation 105ff. microsieve 29, 36f. – metal 36 – polymeric 36 – silicon 26, 36 microsolute transmission 52 migration effects 39, 112 milk – composition 26f., 46 – concentration 5, 46 – consistency 5 – cream stability 27 – fat removal 27 – pH adjustment 5 – powder 55 – processing 3f., 46f. – protein concentrate (MPC) 5 – protein standardization 5, 45f. – skim 3, 5, 27f., 55, 58, 60ff. – taste 8, 27 – total solids 5f., 8 – ultra-high temperature (UHT) 55, 57, 61ff. molecular – distillation 206 – encapsulation 205f. molecular weight cut-offs (MWCOs) 2, 107f., 113, 168, 204f. must – acidiﬁcation 125f. – concentrated (CM) 106, 115 – fermentation 105 – rectiﬁcated concentrated (RCM) 105f., 115 – reduced acidity 120, 123 – sugar reduction 119f.

n nanoﬁltration (NF) 1f. – dairy-process water 66f. – fruit juice concentration 168 – grape must 115f. – milk 50 – rejection coefﬁcients 115 – taste of wine 126 – volatile acidity in wine 124 – whey demineralization 8 neutralization 124f. nitrogenation 22 nonsolvent-induced phase separation (NIPS) technique 234

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o organoleptic 3, 126, 168 Ostwald ripening 152

p partial condensation 168 pasteurization 15, 28, 48 – cold 27, 46 – ultra- 3 permeability – coefﬁcient 227 – gas 223, 225ff. – water-vapor 226 permeate – ﬂow rate 61f. – ﬂux 47f., 56 – ﬂux decline 107f. – recirculation 47, 51 pervaporation (PV) 1f., 21, 168, 227 – aroma recovery 21 – fruit juice 168, 181 pH – adjustment 80, 86f., 123 – milk processing 5 – wine processing 123f. pore size 3f., 26f. – distribution 26, 36 – MF membrane 107f. – uniform 35ff. porosity gradient (GP) 47f. precipitation – polysaccharide 113 – selective 50 – tartrate salts 84 – thermocalcic 6 pressure – capillary penetration 173 – critical penetration 169f. – Laplace 138, 169 – osmotic 2, 146f., 149, 168 – vapor 171 – -variation cycle 57 proteins – bovine serum albumin (BSA) 30, 45, 142f. – denaturation 34 – egg-white 142 – immobilization 211 – immunoglobulins 30 – a-lactalbumin 6, 30, 45 – b-lactalbumin 6, 30, 45, 142, 215 – lactoferrin 30 – repulsion 50

– serum proteins recovery 30 – soybean ﬂour 142 – total concentration 48, 63 – transferrin 30 puriﬁcation 167, 202

r retentate 3 – concentration factor 47 – microﬁltration 7 retention factor 204 reverse osmosis 1f., 81 – fruit juice concentration 177f., 194 – dairy-process water 66 – diaﬁltration 13f. – membrane distillation (MD) 188f. – milk 5f., 50 – must 115 – osmotic distillation (OD) 177ff. – volatile acidity in wine 124 Reynolds number 186 ripening 9, 26, 152

s saltiﬁcation balance 125 salting-out effect 95 shelf life – foods 223, 229, 233 – retention 229 simulated ultraﬁltrate (SMUF) 28 solute rejection rates 52 solution – brine 4 – diffusion 1, 21 – stripping 170, 176, 190 – surface tension 173, 183 spore removal – milk 3f., 46 – skim milk 27 Stacks process 95 starter-culture 5 sterilization cold 28, 30

t tartaric stabilization, see electrodialysis temperature polarization 186, 189 total soluble solids (TSS) 175ff. transmembrane ﬂux 38, 131, 173, 175 transmembrane pressure (TMP) 3, 5, 26, 32, 57, 61, 64 – MF 107ff. – UF 108f. – uniform (UTP) 28, 32f., 46f., 52

Index transmembrane temperature gradient 187, 191 turbidity – fruit juice 181 – permeate 47, 55f., 58 turbulence – micro- 33 – promotion 29, 32

u ultraﬁltration 1f. – cheese production 9, 29, 63 – /electrodialysis (ED) 80 – fruit juice 16, 177ff. – membrane distillation (MD) 188 – milk 5f., 8, 48f. – osmotic distillation (OD) 177ff. – polysaccharide removal 113ff. – powder milk 56 – productivity 107f. – skim milk 60ff. – vinegar 15 – whey demineralization 8 – whey proteins fractionation 52, 64f. – wine clariﬁcation 105ff. UV-light treatment 9

v videomicroscopy 134, 137f. vinegar 14f. viscosity – emulsion 153, 155 – fruit juice 168, 181 – Krieger-Dougherty (KD) equation 155f. – milk 5 volume reduction ratio (VRR) 47, 49f., 56, 61, 64f.

w water – aroma 121 – dairy-process 66 – deoxygenized 22 – treatment 1, 18 – waste- 1, 8, 18, 20, 94 watering down 121 whey – defatted 6f. – demineralization 7f., 30 – processing 6ff. whey protein concentrate (WPC) 4f., 66 whey protein isolate (WPI) 4, 6 whey protein transmission 47, 49, 57ff. wine – acidiﬁcation 125f. – alcohol content reduction 119ff. – clariﬁcation 13, 209 – compounds 14 – concentration factor 107ff. – dealcoholization 12, 14 – fruitiness 123 – pH adjustment 123f. – quality 14, 107, 119, 123 – rejuvenation 13 – stabilization 209 – taste 126, 209 – volatile acidity 123

y yeast – membrane bioreactors 208f., 211 – residues 106 – surplus 10 yoghurt – processing 50 – selective demineralization 50

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Membrane Technology: Volume 3: Membranes for Food Applications