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Studies on the fractionation of β-lactoglobulin from casein whey using ultrafiltration and ion-exchange membrane chromatography
Journal of Membrane Science 275 (2006) 141–150
Studies on the fractionation of -lactoglobulin from casein whey using ultrafiltration and ion-exchange membrane chromatography Sangita Bhattacharjee, Chiranjib Bhattacharjee ∗ , Siddhartha Datta Department of Chemical Engineering, Jadavpur University, Kolkata 700032, India Received 14 February 2005; received in revised form 31 August 2005; accepted 8 September 2005 Available online 10 October 2005
Abstract The objective of this work is to obtain a relative separation of -lactoglobulin (-LG) from whey protein concentrate by fractionation of protein using two-stage ultrafiltration (UF) with 30 and 10 kg mol−1 molecular weight cut-off (MWCO) flat-disk membranes in stirred rotating disk module followed by ion-exchange membrane chromatography (IEMC) using VivapureTM Q Mini-H column. Prior to UF, centrifugation, microfiltration and a four-stage discontinuous diafiltration (DD) were carried out to obtain whey protein concentrate from raw casein whey. Centrifugation and microfiltration were carried out to remove the major foulants like colloidal matters, suspended casein particles and lipid. DD using 5 kg mol−1 polyethersulphone (PES) membrane with volume concentration ratio (VCR) 2 in each stage was employed with an objective to enrich the whey proteins by removing most of lactose, minerals and salts. The effects of various parameters, like solution pH, trans-membrane pressure (TMP), stirrer speed and membrane rotation speed on UF flux and rejection were studied thoroughly. A 36% higher flux was obtained with a feed pH of 2.8 compared to that at pH of 5.6 at a fixed TMP of 5.0 × 105 Pa which has been explained with the help of the effect of pH on monomer–dimer equilibrium of -LG and prevailing protein charge. A 33% enhancement of flux was observed at 1 min with membrane disc rotating at 300 rpm compared to the corresponding value of stationary membrane due to reduction in concentration polarization, the main limiting phenomenon for flux decline. A suitable loading buffer pH was investigated during IEMC runs so as to get a facilitated transport of -LG over ␣-LA through the strong anion exchanger membrane. An 87.6% purity of -LG (on total protein basis) was obtained in the filtrate of ion-exchange membrane chromatography. © 2005 Elsevier B.V. All rights reserved. Keywords: Ultrafiltration; Whey protein; Flux; Rejection; Isoelectric point; Rotating disk membrane module; Ion-exchange membrane chromatography
1. Introduction Ultrafiltration (UF) has been an accepted and well-practised processing operation in the dairy industry. The two largest applications are the fractionation of cheese whey and the preconcentration of milk for cheese manufacture. In cheese indusAbbreviations: BOD, biological oxygen demand (mg L−1 ); COD, chemical oxygen demand (mg L−1 ); DD, discontinuous diafiltration; HPLC, high performance liquid chromatography; HPTFF, high performance tangential flow filtration; IEMC, ion-exchange membrane chromatography; IEP, isoelectric point; J, volumetric flux (L m−2 h−1 ); MF, microfiltration; MWCO, molecular weight cut-off; nm , membrane rotation speed (rpm); ns , stirrer speed (rpm); PES, polyethersulphone; TMP, transmembrane pressure (Pa); UF, ultrafiltration; VCR, volume concentration ratio; WPC, whey protein concentrate; ␣-LA, ␣-Lactalbumin; -LG, -Lactoglobulin ∗ Corresponding author. Tel.: +91 92316 92975; fax: +91 33 2414 6378. E-mail address: [email protected] (C. Bhattacharjee). 0376-7388/$ – see front matter © 2005 Elsevier B.V. All rights reserved. doi:10.1016/j.memsci.2005.09.013
try, almost 10 kg of milk gives 1–2 kg of cheese and 8–9 kg of liquid whey, depending on the quality of the milk. Low solid content and very unfavorable lactose:protein ratio make the utilization of whey a difficult proposition. Liquid whey can be ultrafiltered to produce whey protein concentrates (WPCs) which may be fractionated into individual whey proteins [1,2]. One of the important applications of milk is the casein preparation, which is achieved either by rennet coagulation or by isoelectric precipitation with an acid. The acid may be either a mineral acid or lactic acid. The composition of whey proteins in casein whey, their molecular weights, isoelectric points are tabulated in Table 1 [3]. The disposal of casein whey, having a biological oxygen demand (BOD) value of about 35,000–60,000 mg L−1 and chemical oxygen demand (COD) value of 80,000–100,000 mg L−1 as sewage, cause severe environmental pollution problem. Moreover, the whey protein fractions containing a wide array of proteins namely -lactoglobulin
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Table 1 Major protein composition in bovine (acid) whey
-Lactoglobulins ␣-Lactalbumin Serum albumin IgG, IgA, IgM Lactoperoxidase Lactoferrin
Concentration (kg m−3 )
Isoelectric point (pI)
Molecular weight (kg mol−1 )
3–4 1.2–1.5 0.3–0.6 0.6–0.9 0.06 0.05
5.2–5.4 4.2 4.9–5.1 5.8–7.3 9.6 8
18.3 14.2 66.0 150.0–900.0 78.0 78.0
From Ref. [3].
(-LG), ␣-lactalbumin (␣-LA), bovine serum albumin (BSA), lactoferrin, lactoperoxidase, glycomacropeptide, etc., some of which could have important end uses, are getting disposed off as sewage. Even in WPCs and individual whey protein concentrates (IWPCs), the unique nutritional, therapeutic and functional characteristics of the individual whey proteins are largely unrealized due to interactions between components and degradation during processing. There, thus has been considerable commercial interest in the production of individual whey proteins with well characterized functional and biological properties by a process which will not denature, but retain its nutritional and other properties. Extensive researches have been carried out during last few years keeping this perspective in view [4,5]. -Lactoglobulin (-LG) is a better foam stabilizer than the other whey proteins and can be used in the production of confectionery, whereas ␣-lactalbumin (␣-LA) can be used in infant formula and as a neutraceutical because of its high tryptophan content. It also provides enhanced whippability in meringue like formulations. In addition, ␣-lactalbumin has strong affinity for glycosylated receptors on the surface of oocytes and spermatozoids and may, thus have strong potential as a contraceptive agent [6]. Extensive studies have been carried out by the researchers to demonstrate the feasibility of membrane systems for the separations of proteins with very similar molecular sizes. Van Reis et al. [7] used high performance tangential flow filtration (HPTFF) to separate bovine serum albumin (BSA) from an antigen binding fragment of a monoclonal antibody (Fab), achieving more than 900-fold purification and 90% yield of the BSA. In a previous work carried out by Zhang [8] regarding the study on the affinity cross-flow filtration process for separation of BSA from IgG molecules, the target protein BSA was recovered in more than 95% purity. van Eijndhoven et al. [9] developed a membrane system for the separation of bovine serum albumin (BSA) from hemoglobin, two proteins with essentially identical molecular weight, with more than 100-fold purification and nearly 70% yield. More recently, Cheang and Zydney [10] were able to obtain 100-fold purification and greater than 90% recovery of -LG from a binary mixture with ␣-LA. All these experimental data were obtained with model system consisting of an artificial mixture of two previously purified proteins. Experimental studies with complex multicomponent feed streams are much more limited, and the overall performance of these systems is much less impressive. With an objective to purify ␣-LA from acid casein whey, Muller et al. [11] used a combined ultrafiltration–diafiltration process and an ␣-LA purity of
only 50% was obtained in the final permeate. Bottomley [12] described a two-stage membrane process for the purification of ␣-LA from cheddar whey, but the final product still contained nearly 25% -LG. Lucas et al. [13] demonstrated the extraction of ␣-LA selectively from acid casein whey protein concentrate (WPC) at pH 7 by limiting -LG transmission by chemically modifying the inorganic membrane by a polyethyleneimine coating bearing positive charges. Some studies of proteins fractionation are based on the physico-chemical environment and the charge of the solutes. Mehra and Donnelly [14] showed the possibility of fractionating clarified or dialysed whey with cellulose membranes. The retentate was enriched with the high molecular weight proteins (immunoglobulin, lactoferrin, bovine serum albumin) and the permeate was enriched with the low molecular weight proteins (-lactoglobulin, ␣-lactalbumin). Ion exchange membrane chromatography (IEMC) [15,16] is a high resolution separation technique most suited for protein purification protocols. The separation is based on the reversible electrostatic interaction between a charged protein molecule and the oppositely charged chromatographic membrane (ion exchanger). Amino acids, for example, have both acidic and basic properties and these charge properties have a definite role on the reactivity of proteins to ion-exchange media. An ionexchanger consists of an insoluble matrix, such as cellulose, silica or often a form of polymer, styrene divinylbenzene to which charged groups have been covalently bound. The charged groups are associated with mobile counter-ions. The counterions may be reversibly exchanged with other ions of the same charge. The protein must displace the counter ions and therefore bind to the exchanger on the membrane. Before proceeding to choose the right IEMC column, one must consider whether the target protein would be bound as polyanion or polycation. As proteins are usually handled in buffer media, one is usually interested in their isoelectric points (IEP). The protein is positively charged below IEP and it will bind to any cationic exchange membrane. At a pH higher than the IEP of the protein, the target protein will be negatively charged and bind to anionic exchange membrane. Desorption of biomolecules from the ion-exchange membranes begins after increasing the ionic strength or changing the pH of the elution buffer. The work in this paper has been undertaken with an objective to examine the relative separation of -LG from whey protein concentrate using two-stage ultrafiltration with 30 and 10 kg mol−1 MWCO membranes in a stirred rotating disc membrane module, followed by ion-exchange membrane chromatography. It is reported [17] that at 20 ◦ C, bovine -LG forms a dimer at neutral pH, while the monomeric native state is stable at pH values below 3 and at low ionic strength in absence of salts. To study the overall performance at two different pH, one above and the other below the IEP of casein (IEP of casein being 4.6 [5]), experiments were carried at two different pH, i.e. 2.8 and 5.6. In each stage of ultrafiltration, the effects of stirrer speed, membrane disc rotation, transmembrane pressures (TMP) and solution pH on permeate flux and rejection were thoroughly investigated. The results obtained were compared in terms of the concentration ratio of -LG to ␣-LA and separation factor, in addition to other variables like flux, rejection, etc. One thing
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that needs to be mentioned is that, in IEMC spin column, the driving force is centrifugal which limits its use and scale-up in larger industrial version. But here, the main emphasis is to show the feasibility of the concept, i.e. possibility of obtaining a relative separation between ␣-LA and -LG using anion exchanger membrane, rather than to suggest a process flow-sheet which is directly applicable in industry. 2. Materials and methods 2.1. Materials Experiments were performed using casein whey obtained from local sweet meat industries situated in-and-around Kolkata, India. The pH of the raw casein whey was found to vary from 3 to 4, depending upon the quantity of excess acid present in whey resulting from acid caseination. In most of the cases, the sweet-meat industries use hydrochloric acid or its equivalent, for casein precipitation. Several standard proteins like ␣-lactalbumin (␣-LA), lactoglobulin (-LG), bovine serum albumin (BSA), lactoferrin (LF) and immunoglobulin (Ig) were purchased from either E.Merck (Mumbai, India) or from Sigma–Aldrich (St. Louis, USA) for forming standard solutions, for characterization of high performance liquid chromatography (HPLC) readings. All other chemicals (required for membrane cleaning, storage, and HPLC measurements) were purchased locally, mostly from E.Merck (Mumbai, India). The deionized water used in all the experiments was obtained from Arium 611DI ultrapure water system (make: Sartorius AG, G¨ottingen, Germany). The feed to
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this Arium 611DI was taken from usual laboratory distillation unit. 2.2. Membrane and module Ultrafiltration of pretreated casein whey was carried out batch wise in stirred rotating disc module (capable of being used as fixed disc module also) using polyether sulfone (PES) membrane. The module made of SS316, was manufactured by Gurpreet Engineering Works, Kanpur, UP (India) as per specified design. The module was equipped with two motors (Fig. 1) with speed-controllers to provide rotation of the stirrer and membrane housing. The module has the facility to rotate the membrane and the stirrer in opposite direction to provide maximum shear in the vicinity of the membrane. Digital tachometer was used to measure the rotational speed of both the membrane and the stirrer. Adequate mechanical sealing mechanism was provided to prevent leakage from the rotating membrane assembly. The magnetic drive stirrer mechanism prevents any possibility of leakage from the top stirrer. The complete schematic diagram of the rotating disk module setup was given in Fig. 1. The PES membranes (flat disk of 76 mm diameter) of 5, 10 and 30 kg mol−1 molecular weight cut-off (MWCO) were imported from Millipore Corporation, Bedford, MA (USA) through its Indian counterpart (Millipore India Limited). The 5 kg mol−1 membrane was used for the DD operation, which was also carried out in the same stirred rotating disk module, operated under a fixed set of operating conditions. The flat sheet membrane, operable in pH range of 1–14, has an actual diameter of 76 mm whereas the effective diameter was 56 mm. The membrane is
Fig. 1. Schematic diagram of rotating disc membrane module setup.
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Table 2 Preparation of loading buffer (of volume 100 mL) of different pH
2.3. Pretreatment of casein whey
Loading buffer pH
Amount of 0.1 M citric acid solution (mL)
Amount of 0.1 M trisodium citrate solution (mL)
4.6 4.8 5.0
44.5 40.0 35.0
55.5 60.0 65.0
In order to prevent any possibility of membrane fouling, the suspended casein particles and colloidal matters (mostly fat) were removed by centrifugation followed by microfiltration. Centrifugation was carried out in a research centrifuge model TC 4100D (make: Remi, Mumbai, India) with a speed of 12,500 rpm giving a RCF (relative centrifugal force) of 16,000 g for a period of 30 min. After centrifugation, the sample was subjected to microfiltration (MF) using “all glass vacuum filtration unit” (make: Sartorius AG, G¨ottingen, Germany), fitted with oilfree portable vacuum pump (Sartorius AG, G¨ottingen, Germany, model ROC 300 with moisture trap), with polyether sulfone (PES) membrane (47 mm diameter, pore size 0.45 m, Sartorius AG, G¨ottingen, Germany) being used as filter media. PES membranes exhibit no hydrophobic or hydrophilic interactions and are usually preferred for their low fouling characteristics, broad pH range and durability. The permeate from the microfiltration still contains lactose and different minerals, which needs to be removed to obtain, so called whey protein concentrate (WPC). The removal of lactose is very much essential as it may impair the downstream separation operation. Lactose and all other mineral salts were removed by a four-stage discontinuous diafiltration (DD) with 5 kg mol−1 PES membrane, the detail procedure of which will be mentioned later. The final permeate from the DD was adjusted for pH by adding calculated amount of hydrochloric acid or sodium hydroxide, as required to produce the feed for the subsequent UF run at two desired pHs, i.e 2.8 and 5.6.
compatible both in alcohol and aqueous medium and can be sterilized by immersion in 3% formalin or 5% hydrogen peroxide or 0.1% peracetic acid for 24 h. The retentate obtained after carrying UF using 10 kg mol−1 membrane (which mainly contains only -LG and ␣-LA) was subjected to ion-exchange membrane chromatography (IEMC) using VivapureTM Q Mini-H column (manufactured by Vivascience AG, Hannover, Germany; supplied by Sartorius Biotech Pvt. Limited, Bangalore, India) which has a strong basic anion exchanger with quaternary ammonium ion [R CH2 N+ (CH3 )3 Cl− ] immobilized on regenerated cellulose acetate substrate. This membrane was housed in a spin column with centrifugal action being used as the driving force. For equilibration of the membrane, 25 mM Tris/HCl, 25 mM NaCl buffer pH 8.0 was supplied with VivapureTM Q Mini-H column. Loading buffer of different pH was prepared by mixing citric acid solution and trisodium citrate solution. Citric acid monohydrate and trisodium citrate dihydrate were purchased from E-Merck, Germany. During experimentation, three different levels of loading buffer pH were selected to study optimum recovery of -LG. The levels were 4.6, 4.8 and 5.0, and the method of preparation of these loading buffer were mentioned in Table 2. Complete process flow diagram showing all the steps carried out during this analysis has been shown in Fig. 2.
2.4. Membrane compaction and water run Prior to use in experiments, each membrane was subjected to compaction for about 1 h with ultra-pure water at a pres-
Fig. 2. Schematic diagram of the proposed process flow diagram.
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sure of 9 × 105 Pa, higher than the highest operating pressure used in this study, to prevent any possibility of change of membrane hydraulic resistance during ultrafiltration. Once the water flux becomes steady with no further decrease, it was concluded that full compaction of the membrane has taken place. After compaction, membrane hydraulic resistance (Rm ) was determined based on water runs at different transmembrane pressures (TMPs) of (4–7) × 105 Pa. Membrane was rinsed thoroughly with distilled water after every run with casein whey so as to remove any deposited fouling layer, which was followed by water runs to determine the extent of fouling. The drops in water fluxes obtained from such studies were found to be within 2% of initial water flux, thus showing minimum fouling resulting from the proposed separation scheme. 2.5. Methodology The centrifuged and microfiltered whey was batch diafiltered four times using 5 kg mol−1 polyether sulphone (PES) membrane in the above module at 500 rpm stirrer speed, 300 rpm membrane rotation speed and TMP of 4 × 105 Pa with an initial volume 350 mL. During first stage of DD, the operation was carried out until a volume concentration ratio (VCR) of 2 was achieved. The required VCR was attained in approximately 1.2 h, after which 175 mL of deionised water was added and second stage diafiltration was carried out. This procedure was continued in all the four-stages of diafiltration operation. Ultrafiltration experiments were carried out batch wise in rotating disc membrane module at different operating conditions starting each time with a pretreated and diafiltered feed volume of 350 mL (175 mL of deionised water was added to final DD retentate to makeup the volume up to 350 mL). As per the objectives of the experiments, the solution pH was varied at 2.8 and 5.6, and the TMP was fixed at (4–7) × 105 Pa. The stirrer speed (ns ) was set as 0 (fixed stirrer), 500 and 1000 rpm, whereas the membrane rotation speed (nm ) was varied at 0 (fixed membrane), 300 and 600 rpm. The stirrer speed and membrane rotation speed were controlled through speed controllers, regulating the power supply to the respective motors. The stirrer speed and the membrane rotational speed were checked with the help of a digital tachometer working on stroboscopic principle during experiment. The experiments were designed in such a manner so as to get the nature of variation of flux and rejection with each of the independent variables as mentioned above. The permeate of 30 kg mol−1 membrane was again subjected to UF in rotating disc module using 10 kg mol−1 membrane and all the experimental runs were taken at different values of TMPs, stirrer speeds, membrane rotation speeds and pHs, as mentioned earlier. After each experiment, the module was disassembled first, the membrane disc was thoroughly rinsed for 20 min under running water in laboratory wash basin, soaked in 1 L of 0.5 mM NaOCl in 0.5 M NaOH for 30 min in a beaker, subsequently soaked in deionised water for 5 h and then again thoroughly washed for 20 min. In each case, the water flux was found to regain by more than 98% of its original value, suggesting the cause of flux decline to be either osmotic pressure limited or due to reversible fouling layer.
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The two types of retentate of 10 kg mol−1 membrane, resulting from initial feed solutions at two different pHs (2.8 and 5.6) were then fed to ion-exchange membrane chromatography unit. Ion-exchange membrane chromatography (IEMC) is a high resolution separation technique most suited for protein purification protocols. The separation is based on the reversible electrostatic interaction between a charged chromatographic membrane (ion exchanger) and the target protein molecule. In this work, VivapureTM Q Mini-H column manufactured by Vivascience AG (Hannover, Germany), supplied by Sartorius Biotech Pvt. Limited (Bangalore, India) was used for this purpose. The unit comprises of a strong basic anion exchanger, in the form of quaternary ammonium ion (R CH2 N+ (CH3 )3 Cl− ), with a membrane of bed volume 240 L, membrane surface area 7.48 × 10−4 m2 and amount of ionic groups to be 145–218 equiv. mL−1 for monovalent ions, and the membrane supported on stabilized regenerated cellulose acetate matrix. Following steps were involved in the proposed ion-exchange membrane chromatographic separation: (1) Equilibration of the VivapureTM Q Mini-H spin column with 400 L of 25 mM Tris/HCl and 25 mM NaCl buffer at pH 8.0, and subsequent spinning at 2000 g for 5 min in fixed angle centrifuge. (2) An amount of 250 L loading buffer of desired pH, prepared from citric acid monohydrate and trisodium citrate dihydrate (see Table 2) was added to 50 L of sample and vigorously mixed. (3) An amount of 300 L diluted sample was loaded onto VivapureTM Q Mini-H column and rotated at 2000 g for 5 min. Flow-through-fraction sample contained enriched LG. (4) An amount of 50 L of above mentioned buffer of desired pH was loaded again to ensure complete recovery of unbound -LG. (5) The bound protein ␣-LA was recovered by loading 300 L of 25 mM Tris/HCl, 1 M NaCl buffer pH 8.0. The pH of loading buffer was selected at a pH 4.6, 4.8 and 5.0 with an understanding that ␣-LA (pI 4.5–4.8), would be negatively charged at this pH and would bind to the membrane whereas -LG being positively charged (pI being 5.2–5.4) would pass through the membrane surface. The filtrate thus obtained from IEMC was analyzed further for individual protein concentration measurement. Another important factor in this analysis was to study the repeatability/reproducibility of the results obtained through the proposed separation train. For this, all the experiments were repeated under the same operating conditions and in the same modules at least thrice (in some cases, even six to seven times), to assess the reproducibility of the results. In all the cases, deviations between the various runs were found to be minimal, mostly remaining within ±5%. In some instance, greater deviations of one or two points were indeed obtained, which may be due to the experimental error. The repeatability of the results have been indicated in various figures in the form of error bars, where those one or two points showing greater than 10% deviation have been
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considered as off-shoot points and were not included in the calculation of error bars. In some of the figures, the error bars were not shown to preserve the clarity and usability of the figures. 2.6. Analysis The total solid concentration (TS) was determined gravimetrically after drying the samples in an incubator at 105 ◦ C and measuring the weight of the residue. Total protein concentration (TPC) was measured using a Hitachi dual beam UV–vis spectrophotometer (Model No. U-2800, with UV solution software) by the principle of Bradford protein assay at 590 nm. The assay is based on the observation that the maximum absorbance for an acidic solution of Coomassie Brilliant Blue G-250 shifts from 465 to 590 nm when binding to protein occurs [18]. Regarding the individual protein concentration, the focus in this study was mainly to measure ␣-LA and -LG. High performance liquid chromatography (HPLC) was used in this study to measure individual protein concentration. Water® (model 600 S) (imported from Waters Associates, Milford, MA, USA) HPLC with Shodex size exclusion column (8 mm × 300 mm, MW range: 100–150,000 Da) was used for the experimentation purpose. The HPLC was operated with phosphate buffer (pH 6.8) and the flow-rate was fixed at 1 mL min−1 . Disodium hydrogen phosphate (Na2 HPO4 ) and sodium dihydrogen phosphate (NaH2 PO4 ) solutions were prepared each of 0.1 M concentration and 0.15 M sodium chloride (NaCl) was added to these mixtures. These phosphate buffers were used in the ratio of 40/60 and good separation was observed. Detection was done at UV-210 nm. The elution times in most of the cases were within 15 min. In this study, individual protein concentrations were also determined using reversed phase-high performance liquid chromatography (RP-HPLC) according to a modification of the method suggested by Resmini et al. [19] with a silica based wide pore C18 column (15 m, 3.9 mm × 300 mm) in the same Water® Gradient HPLC. In this case, detection was done at UV280 nm and the flow rate fixed at 1 mL min−1 . The mobile phase was a binary system of water with 0.1% of trifluoro acetic acid (TFA) (buffer A) and acetonitrile with 0.1% of TFA (buffer B). The gradient condition was consisted of a two-step linear binary gradient: buffer A/buffer B = 70/30–35/65 (vol.%), gradient time of 15 min. The second method of determination produced a little bit better separation of the protein concerned (i.e. ␣-LA and -LG), though in this study, most of the results were obtained using the first method of size exclusion chromatography. In this study, during interpretation of IEMC results, a parameter called ‘separation factor’ (S) defined by: S=
(-LG : ␣-LA)filtrate (-LG : ␣-LA)feed
(1)
has been used, where, subscript filtrate and feed corresponds to those in ion-exchange chromatography membrane. Separation factor was plotted against loading buffer pH used in VivapureTM Q Mini-H column.
Fig. 3. Flux in four-stage discontinuous diafiltration operation.
3. Result and discussion After membrane compaction, series of water runs were taken to evaluate membrane hydraulic resistance Rm which was found to be (8.264 × 1012 ± 3.7 × 1011 ) m−1 for 30 kg mol−1 membrane and (1.998 × 1013 ± 9.6 × 1011 ) m−1 for 10 kg mol−1 membrane. The higher value of membrane hydraulic resistance for 10 kg mol−1 membrane compared to 30 kg mol−1 membrane is due to reduced pore size and compact nature of the membrane. The variation of flux with cumulative time during four-stage diafiltration with 5 kg mol−1 PES membrane was depicted in Fig. 3. It can be observed from Table 3 that only the pretreatment step alone, consisting of centrifugation, microfiltration and diafiltration, produced an improvement of %-purity of -LG with respect to total protein from 45.2% to 52.4%. The most important is the reduction of total solid (TS) in the pretreatment stage, from 645 kg m−3 in the raw casein whey to 13.89 kg m−3 in the pretreated casein whey, which accounts for 97.8% reduction in total solids. The major reduction in the total solid could be attributed to diafiltration step, which is very much suitable for purification of high molecular weight fraction. Due to reduction in total solid, the whey protein fraction on dry basis (with respect to total solid) has improved significantly, from 0.54% in the raw casein whey to 23.2% in pretreated casein whey. It may be observed from Fig. 3 that during the first stage of diafiltration, the flux has reduced from 140.6 to 11.2 L m−2 h−1 after about 1.2 h when a VCR of 2 was achieved. The protein fraction could even be increased much more by continuing the diafiltration for additional stages and using a little higher VCR. But in this study, that was not our objective, the emphasis was mainly to study the relative separation between ␣-LA and -LG using IEMC and the effect of feed pH on the whole separation train. Detailed results of the suggested two-stage UF employing 30 and 10 kg mol−1 membrane are shown in Table 3 under constant operating condition (pH 2.8, TMP = 4 × 105 Pa, ns = 0 rpm and nm = 300 rpm). Along with the UF results, this table also highlights the findings from IEMC. Table 3 gives the concentration of ␣-LA, -LG and TPC in pretreated casein whey, as
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Table 3 Compositions of pretreated casein whey, UF permeate (30 and 10 kg mol−1 ) and UF retentate (10 kg mol−1 ) at different pH, under fixed operating conditions (TMP = 4 × 105 Pa, ns = 500 rpm and nm = 0 rpm) and results of IEMC with UF-10 kg mol−1 retentate Item
Concentration in raw casein whey (kg m−3 )
Concentration in pretreated casein whey (kg m−3 )
pH
␣-LA
1.26 ± 0.0466
0.93 ± 0.032
-LG
3.46 ± 0.13
3.23 ± 0.12
TPC
7.66 ± 0.29
6.16 ± 0.23
TS
645 ± 25
2.8 5.6 2.8 5.6 2.8 5.6 2.8 5.6 2.8 5.6
13.89 ± 0.5
%-Purity of -LG 45.2 ± 2.1 52.4 ± 2.6 with respect to TPC %-Purity of -LG 0.54 ± 0.032 23.2 ± 1.4 with respect to TS a These concentrations refer to those in the diluted samples.
UF permeate (30 kg mol−1 )
2.8 5.6
well as those in UF-30 kg mol−1 permeate, UF-10 kg mol−1 permeate and UF-10 kg mol−1 retentate. In addition to the protein concentrations, Table 3 also presents total solid (TS) at different stages of filtration, along with %-purity of -LG on total protein, as well as on total solid basis. Table 3 also highlights the effect of feed pH on concentrations of different proteins, as well as on TS and TPC. It could be observed that all the permeate concentrations were greater at different stages at pH 2.8, compared to that at 5.6. It is already reported [17], that bovine -lactoglobulin assumes a monomeric native structure at pH 3 in the absence of salt. This fact is also supported by another article [20] wherein it is also mentioned that at pH 6.2 -LG exists as dimer. Hence greater rejection of -LG at pH 5.6 is attributed due to forward shift of monomer–dimer equilibrium, thereby causing formation of major proportion of -LG which is retained by the membranes (10 and 30 kg mol−1 ) resulting in lesser -LG concentration in UF permeate as well as reduced TS and TPC values at all subsequent stages. The effects of transmembrane pressure (TMP) on flux at pH 2.8 have been shown in Fig. 4 with stationary membrane and stirrer rotating at 500 rpm. Several models were tested to find which fitted the transient variation of flux with respect to time under these conditions of operations. Statistical analysis proved that the relationship between flux and time could be best described by the model given by Sheppard and Thomas [21] based on minimum value of standard error and maximum value of correlation coefficient: J = J1 t −b
Concentration (kg m−3 )
0.81 0.80 2.73 2.29 3.87 3.51 11.50 11.15 70.6 65.1
± ± ± ± ± ± ± ± ± ±
0.028 0.025 0.1 0.09 0.15 0.14 0.41 0.39 4.2 3.9
23.7 ± 1.64 20.5 ± 1.22
UF permeate (10 kg mol−1 ) 0.32 0.28 0.43 0.21 0.83 0.71 7.80 7.66 52. 4 30.4
± ± ± ± ± ± ± ± ± ±
0.011 0.009 0.013 0.007 0.03 0.024 0.31 0.31 3.1 1.9
5. 6 ± 0.38 2. 8 ± 0.18
UF retentate (10 kg mol−1 ) 1.02 0.99 3.71 3.05 5.11 4.55 15.20 14.64 72.7 67.0
± ± ± ± ± ± ± ± ± ±
0.035 0.034 0.14 0.11 0.19 0.17 0.61 0.54 4.5 4.6
24.4 ± 1.46 20.8 ± 1.45
IEMCa permeate (buffer pH 5.0) 0.031 0.028 0.539 0.437 0.615 0.502 2. 050 1.975 87.6 87.0
± ± ± ± ± ± ± ± ± ±
0.001 0.0008 0.02 0.017 0.024 0.018 0.077 0.098 5.25 5.1
26.3 ± 1.55 22.1 ± 1.32
sion parameters, J1 and b. Standard error was found to be 2.1417 and the correlation coefficient being 0.9979. Similarly, J1 and b was computed at other operating conditions. For all the runs, high values of correlation coefficients (>0.95) signifies good fitting of flux versus time data by Sheppard and Thomas model. The flux was observed to decline with time, which is obvious in any pressure-driven membrane process like UF; a 67% reduction of permeate flux was observed within a time span of 20 min at a TMP of 4 × 105 Pa whereas a reduction of 74.5% was observed within the same time period at a TMP of 7 × 105 Pa. This is probably due to the fouling caused by the whey proteins mainly BSA, IgG and partly ␣-LA and -LG onto the membrane surface because of possible formation of gel layer [2]. From the inset graph of Fig. 4, a 36% higher flux was obtained with a feed pH of 2.8 compared to the flux obtained with feed
(2)
where J is the volumetric flux in L m−2 h−1 measured at time t in seconds. The numerical values of J1 and b corresponding to the 4 × 105 Pa TMP and the specified values of nm and ns as shown in figure, have been estimated as 422.99 ± 11.42 and 0.3107 ± 0.0065, respectively, for the aforementioned units of flux and time. Nonlinear regression technique employing method of least-square was used to find the values of the regres-
Fig. 4. Effect of transmembrane pressures on whey flux at constant ns and constant nm using 30 kg mol−1 membrane. Effect of feed solution pH on flux vs. TMP plot at specified TMP, ns and nm .
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pH of 5.6 at a TMP of 5 × 105 Pa. This trend is valid at other operating conditions also. -LG exists in a dimmer–monomer equilibrium between pH 3.5 and 6.5 and as slight expanded monomer below pH 3.5 and above 7.5. Detail effects of pH on -LG and the resulting monomer–octomer–dimer equilibrium have been discussed elsewhere [17,22,23]. It has been already mentioned that, after four stages of discontinuous diafiltration of raw casein whey, the TS value in the pretreated casein whey has been reduced by 97.8%. Hence, the presence of minerals and salts may be considered to be almost negligible in the pretreated casein whey. Therefore, it may be stated that the effect of pH on protein charge was predominating at this low ionic strength. Apart from the presence of major fraction of -LG in monomeric form at pH 2.8, the less dense nature of deposited protein layer onto membrane surface at this pH is another reason for higher flux compared to that at pH 5.6. At pH 2.8, the magnitude of charge on -LG is more compared to that at 5.6 (IEP being 5.2–5.4). Due to stronger [24] mutual electrostatic repulsion among charged -LG molecules at pH 2.8, the deposited protein layer onto membrane surface is less dense. As a result, the concentration boundary layer is less dense causing the effect of polarized layer resistance to a lower value. Fig. 5 shows the variation of permeate flux as a function of time at different membrane rotation speeds under constant pH, ns and TMP. The decline of flux with time is obvious for any pressure-driven membrane processes like UF, which is due to concentration polarization and reversible or irreversible fouling. The flux enhancement with increase in membrane rotation speed could be observed from the figure, which resulted from reduced concentration polarization for a rotating membrane disc. It was observed that flux increased by nearly 33% at 1 min for a rotating membrane (300 rpm) compared to the corresponding
Fig. 5. Effect of membrane rotation on flux vs. time plot during UF of whey at a particular pH using 30 kg mol−1 membrane at constant TMP, ns . (Inset) Comparison of effect of stirrer speed on flux between 30 and 10 kg mol−1 membranes at a specified time at constant TMP and pH.
value for stationary membrane. The concentration polarization and the resulting fouling for stationary membrane could become very severe as the UF process goes on, which reduced the flux very drastically. The flux-decline between 30 s and 20 min of operation for stationary membrane was 67% whereas that for rotating membrane at 300 rpm, it was 46%. Effect of stirrer speed on flux at a particular time (5 min, J5 ) for both the 30 and 10 kg mol−1 membranes have been shown in the inset of Fig. 5, which clearly shows the flux enhancement with stirrer speed, provided all other operating variables remain constant. Increased stirrer speed creates more turbulence near the vicinity of the membranes which subsequently reduces the polarized layer resistance and accordingly increases the flux. Another important fact that can be noticed from the figure is the significantly higher flux of 30 kg mol−1 membrane in comparison to 10 kg mol−1 membrane under same condition of operations. This is expected as 10 kg mol−1 membranes are much more compact with less average pore size, thus giving reduced flux in comparison to 30 kg mol−1 membrane. Variations of observed rejection (R) with TMP for different solution pH under fixed nm and ns are shown in Fig. 6 for UF operations with 30 kg mol−1 membrane. The rejection was higher for pH 5.6, compared to the corresponding results at pH 2.8, the reasons of which have already been discussed earlier. Under a fixed set of operating conditions, the average rejections were found to increase with TMP, which is in somewhat contradiction with the observation in many cases, where apparent rejection has been found to decrease with TMP. Higher TMP would result in higher convective flow and accordingly more
Fig. 6. Effect of TMP on rejection during UF using 30 kg mol−1 membrane at two different pH at constant stirrer speed and with stationary membrane. (Inset) Effect of membrane rotation speed on rejection vs. TMP at constant stirrer speed.
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Fig. 7. Plot of value of the ratio -LG to ␣-LA at four-stage of processing, i.e. pretreatment, first stage UF (30 kg mol−1 ), second stage UF (10 kg mol−1 ) and final stage, i.e. ion-exchange-membrane process.
solute would be transported towards the membrane surface. This would results in higher thickness of the polarized layer as well as high membrane surface concentration. In this case, the increase of rejection with TMP may be explained as follows. Higher pressure has caused more compaction of the polarized layer; the layer as a result has acted as a secondary membrane, thus increasing the rejection due to steric hindrance. Effect of membrane rotation on rejection is shown in the inset of Fig. 6. Membrane rotation was found to decrease the observed rejection, which could be attributed due to reduced polarized layer thickness, thus reducing the secondary membrane effect and increasing solute transport through the membrane. Fig. 7 depicts the pattern of improvement in the ratio of (LG:␣-LA) at different stages and at two different pHs. The first stage corresponds to the pretreated casein whey, second stage corresponds to UF permeate from 30 kg mol−1 membrane and third stage the UF retentate from 10 kg mol−1 membrane with the fourth stage being the permeate (eluted stream) from IEMC. The average values of the ratio of -LG to ␣-LA at each stage were calculated from Table 3 and reported at two different feed solution pH, i.e.2.8 and 5.6. It may be observed that the value of the above mentioned ratio in the permeate of 30 kg mol−1 membrane at initial feed pH 2.8 was 3.37, whereas the value of this ratio was 2.85 corresponding to feed solution pH of 5.6. The retentate of 10 kg mol−1 membrane corresponding to the permeate of 30 kg mol−1 membrane of initial pH of 2.8 was found to give a higher value (3.64) of the ratio of -LG to ␣-LA compared to corresponding value (3.07) of 10 kg mol−1 retentate obtained from 30 kg mol−1 permeate of initial feed solution having pH of 5.6. The reasons of such findings have been discussed earlier. Finally, the stage 4 corresponds to IEMC, the results of which have been shown for loading buffer pH of 5.0 (for which the highest separation was obtained). The ratio (-LG:␣-LA) has improved substantially during the final stage in the elution phase (permeate) of ion-exchange membrane chromatographic separation process. The ratio corresponding to the initial feed solution pH 2.8 was calculated as 17.15 which was substantially higher
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Fig. 8. Plot of separation factor vs. pH of loading buffer in ion-exchange membrane chromatography (IEMC).
than the magnitude of the ratio (9.64) corresponding to feed solution of pH 5.6. In IEMC, for the case reported in Fig. 7, the pH was adjusted in both cases to 5.0 using loading buffer, but the difference in the ratio (-LG:␣-LA) solely resulted from the difference in -LG content in UF-10 kg mol−1 retentate (which was used as feed to IEMC) corresponding to two different pH values. The variation of separation factor as a function of the loading phase pH in the ion-exchange affinity membrane separation process has been shown in Fig. 8. It may be observed from Fig. 8 that a separation factor of 4.71 was obtained against a loading buffer pH of 5.0 whereas only 1.42 was obtained when the sample was mixed with loading buffer of 4.6. This is possibly due to the fact that at loading buffer pH 5.0, the ␣-LA in sample remains as strongly negatively charged (IEP of ␣-LA being 4.2) and was more likely to be bound to the strong basic anion exchanger membrane whereas almost all the -LG passes through the membrane. In contrast, at loading buffer pH 4.8, the negative charge on ␣-LA is less in comparison to that at pH 5.0; hence a lesser fraction of ␣-LA got bound to the membrane. This is probably the reason why separation factor values decreased rapidly as the pH was decreased from 5.0 to 4.6. 4. Conclusion In the present work, detailed investigations on fractionation of casein whey by performing ultrafiltration in rotating disc membrane module was done with an objective to purify -LG in final product. Several pretreatments like centrifugation, microfiltration and diafiltration were used to remove suspended casein, lipids, lactose and other minerals. Diafiltration was found to be very important and using four-stage diafiltration with VCR 2 in each stage, a large improvement of -LG purity with respect to total solid was observed. In fact, -LG purity increased from 0.536% to 23.2% using these diafiltration steps. After pretreatment, two-stage UF, using 30 and 10 kg mol−1 membrane was used, followed by IEMC. Effects of various independent variables in UF, like feed solution pH, transmembrane pressure, stirrer speed, membrane rotation speed were investigated in detail
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and suitable separation scheme has been formulated. Higher stirrer speed and membrane rotation speed justifiably gave higher permeate flux though their rotational speeds were needed to be optimized. Higher flux and lower rejection were obtained in UF with the membranes, 30 and 10 kg mol−1 , at pH 2.8, compared to that at pH 5.6. This could be explained based on the effect of pH on possible dimerization of -LG between pH ranges from 3.5 to 6.5 and protein charge at this low ionic strength, which has also been reported elsewhere. It was observed through this study that to obtain highest separation of ␣-LA and -LG, it is recommended to keep the solution pH of pretreated whey in the range where -LG would remain the monomer state. This would result in highest separation and at pH 2.8, the ratio -LG/␣-LA was found to be 17.15 compared to 9.64 at pH 5.6. Further, selection of loading buffer pH in IEMC was found to be very important and in this case, a loading buffer pH of 5.0 was found to give best separation of ␣-LA and -LG. This study resulted in good relative separation of ␣-LA and -LG, and -LG purity of nearly 87% on total protein basis was obtained, though the purity was quite low (26%) on total solid (dry) basis. To improve this percentage, the lactose and other minerals, which still remained after four-stage diafiltration steps need to be removed. This requires some further study, but it could be accomplished by increasing the VCR in diafiltration step and continuing for more steps in diafiltration. Acknowledgement All the works in this study has been performed on equipment purchased under AICTE (TAPTEC) project (No. 8021/RID/NPROJ/TAP-6/2002-03) and are the part of the corresponding project work. Accordingly, the contribution of All India Council for Technical Education (AICTE), Government of India is gratefully acknowledged. References [1] K. Scott, Handbook of Industrial Membranes, Elsevier, Oxford, 1995. [2] M. Cheryan, Ultrafiltration and Microfiltration Handbook, Technomic Publications, Lancaster, 1998. [3] L. Pedersen, J. Mollerup, E. Hansen, A. Jungbauer, Whey proteins as a model system for chromatographic separation of proteins, J. Chromatogr. B 790 (2003) 161–173. [4] E.E.G. Rojas, J.S.D.R. Coimbra, Size-exclusion chromatography applied to the purification of whey proteins from the polymeric and saline phases of aqueous two-phase system, Process Biochem. 39 (2004) 1751– 1759.
[5] P. Riechel, T. Weiss, M. Weiss, R. Ulber, H. Buchholz, T. Scheper, Determination of minor whey protein bovine lactoferrin in cheese whey concentrates with capillary electrophoresis, J. Chromatogr. A 817 (1998) 187–193. [6] A.L. Zydney, Protein separations using membrane filtration: new opportunities for whey fractionation, Int. Dairy J. 8 (1998) 243–250. [7] R. van Reis, J.M. Brake, J. Charkoudian, D.B. Burns, A.L. Zydney, High performance tangential flow filtration using charged membranes, J. Membr. Sci. 159 (1999) 133–142. [8] W. Zhang, A study on the affinity cross-flow filtration process, Ph.D. Thesis, University of Twente, Netherlands, 1993. [9] H.C.M. van Eijndhoven, S. Saksena, A.L. Zydney, Protein fractionation using electrostatic interactions in membrane filtration, Biotechnol. Bioeng. 48 (1995) 406–414. [10] B.L. Cheang, A.L. Zydney, Separation of ␣-lactalbumin and lactoglobulin using membrane ultrafiltration, Biotechnol. Bioeng. 83 (2003) 201–209. [11] A. Muller, G. Daufin, B. Chaufer, Ultrafiltration modes of operation for the separation of ␣-Lactalbumin from acid casein whey, J. Membr. Sci. 153 (1999) 9–21. [12] R.C. Bottomley, Process for obtaining concentrates having a high ␣lactalbumin content from whey. US Patent No. 5,008,376 (1991). [13] D. Lucas, M. Rabiller-Baudry, L. Millesime, B. Chaufer, G. Daufin, J. Membr. Sci. 148 (1998) 1–12. [14] R.K. Mehra, J.W. Donnely, Fractionation of whey protein components through a large pore size, hydrophilic, cellulosic membrane, J. Dairy Res. 60 (1993) 89. [15] C. Charcosset, Purification of proteins by membrane chromatography, J. Chem. Technol. Biotechnol. 71 (1998) 95–110. [16] C. Chiu, M. Etzel, Fractionation of lactoperoxidase and lactoferrin from bovine whey using a cation exchange membrane, J. Food Sci. 62 (1998) 996–1000. [17] K. Sakurai, M. Oobatake, Y. Goto, Salt-dependent monomer-dimer equilibrium of bovine -lactoglobulin at pH-3, Protein Sci. 10 (2001) 2325–2335. [18] M.M. Bradford, A rapid and sensitive method for the quantification of microgram quantities of protein utilizing the principle of protein-dye binding, Anal. Biochem. 72 (1976) 248–254. [19] P. Resmini, L. Pelligrino, R. Andreini, F. Prati, Determination of soluble milk serum proteins by reversed-phase high pressure liquid chromatography (HPLC), Sci. E. Technica Latteriocaseria 40 (1989) 7. [20] K. Kuwata, M. Hoshino, V. Forge, S. Era, CA. Batt, Y. Goto, Solution structure and dynamics of bovine beta-globulin A, Protein Sci. 8 (1999) 2541–2545. [21] J.D. Sheppard, D.G. Thomas, in: A.F. Turbak (Ed.), Membranes from Cellulose and Cellulose Derivatives, American Chemical Society, Washington, DC, 1970, p. 121. [22] J.I. Boye, A.A. Ismail, I. Alli, Effect of physico-chemical factors on the secondary structure of -lactoglobulin, J. Dairy Res. 63 (1996) 97–109. [23] H.A. McKenzie, W.H. Sawyer, Effect of pH on -lactoglobulin, Nature 214 (1967) 1101–1104. [24] Y. Mukai, E. Iritani, T. Murase, Effect of protein charge on cake properties in dead-end ultrafiltration of protein solutions, J. Membr. Sci. 137 (1997) 271–275.
Studies on the fractionation of -lactoglobulin from casein whey using ultrafiltration and ion-exchange membrane chromatography Sangita Bhattacharjee, Chiranjib Bhattacharjee ∗ , Siddhartha Datta Department of Chemical Engineering, Jadavpur University, Kolkata 700032, India Received 14 February 2005; received in revised form 31 August 2005; accepted 8 September 2005 Available online 10 October 2005
Abstract The objective of this work is to obtain a relative separation of -lactoglobulin (-LG) from whey protein concentrate by fractionation of protein using two-stage ultrafiltration (UF) with 30 and 10 kg mol−1 molecular weight cut-off (MWCO) flat-disk membranes in stirred rotating disk module followed by ion-exchange membrane chromatography (IEMC) using VivapureTM Q Mini-H column. Prior to UF, centrifugation, microfiltration and a four-stage discontinuous diafiltration (DD) were carried out to obtain whey protein concentrate from raw casein whey. Centrifugation and microfiltration were carried out to remove the major foulants like colloidal matters, suspended casein particles and lipid. DD using 5 kg mol−1 polyethersulphone (PES) membrane with volume concentration ratio (VCR) 2 in each stage was employed with an objective to enrich the whey proteins by removing most of lactose, minerals and salts. The effects of various parameters, like solution pH, trans-membrane pressure (TMP), stirrer speed and membrane rotation speed on UF flux and rejection were studied thoroughly. A 36% higher flux was obtained with a feed pH of 2.8 compared to that at pH of 5.6 at a fixed TMP of 5.0 × 105 Pa which has been explained with the help of the effect of pH on monomer–dimer equilibrium of -LG and prevailing protein charge. A 33% enhancement of flux was observed at 1 min with membrane disc rotating at 300 rpm compared to the corresponding value of stationary membrane due to reduction in concentration polarization, the main limiting phenomenon for flux decline. A suitable loading buffer pH was investigated during IEMC runs so as to get a facilitated transport of -LG over ␣-LA through the strong anion exchanger membrane. An 87.6% purity of -LG (on total protein basis) was obtained in the filtrate of ion-exchange membrane chromatography. © 2005 Elsevier B.V. All rights reserved. Keywords: Ultrafiltration; Whey protein; Flux; Rejection; Isoelectric point; Rotating disk membrane module; Ion-exchange membrane chromatography
1. Introduction Ultrafiltration (UF) has been an accepted and well-practised processing operation in the dairy industry. The two largest applications are the fractionation of cheese whey and the preconcentration of milk for cheese manufacture. In cheese indusAbbreviations: BOD, biological oxygen demand (mg L−1 ); COD, chemical oxygen demand (mg L−1 ); DD, discontinuous diafiltration; HPLC, high performance liquid chromatography; HPTFF, high performance tangential flow filtration; IEMC, ion-exchange membrane chromatography; IEP, isoelectric point; J, volumetric flux (L m−2 h−1 ); MF, microfiltration; MWCO, molecular weight cut-off; nm , membrane rotation speed (rpm); ns , stirrer speed (rpm); PES, polyethersulphone; TMP, transmembrane pressure (Pa); UF, ultrafiltration; VCR, volume concentration ratio; WPC, whey protein concentrate; ␣-LA, ␣-Lactalbumin; -LG, -Lactoglobulin ∗ Corresponding author. Tel.: +91 92316 92975; fax: +91 33 2414 6378. E-mail address: [email protected] (C. Bhattacharjee). 0376-7388/$ – see front matter © 2005 Elsevier B.V. All rights reserved. doi:10.1016/j.memsci.2005.09.013
try, almost 10 kg of milk gives 1–2 kg of cheese and 8–9 kg of liquid whey, depending on the quality of the milk. Low solid content and very unfavorable lactose:protein ratio make the utilization of whey a difficult proposition. Liquid whey can be ultrafiltered to produce whey protein concentrates (WPCs) which may be fractionated into individual whey proteins [1,2]. One of the important applications of milk is the casein preparation, which is achieved either by rennet coagulation or by isoelectric precipitation with an acid. The acid may be either a mineral acid or lactic acid. The composition of whey proteins in casein whey, their molecular weights, isoelectric points are tabulated in Table 1 [3]. The disposal of casein whey, having a biological oxygen demand (BOD) value of about 35,000–60,000 mg L−1 and chemical oxygen demand (COD) value of 80,000–100,000 mg L−1 as sewage, cause severe environmental pollution problem. Moreover, the whey protein fractions containing a wide array of proteins namely -lactoglobulin
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Table 1 Major protein composition in bovine (acid) whey
-Lactoglobulins ␣-Lactalbumin Serum albumin IgG, IgA, IgM Lactoperoxidase Lactoferrin
Concentration (kg m−3 )
Isoelectric point (pI)
Molecular weight (kg mol−1 )
3–4 1.2–1.5 0.3–0.6 0.6–0.9 0.06 0.05
5.2–5.4 4.2 4.9–5.1 5.8–7.3 9.6 8
18.3 14.2 66.0 150.0–900.0 78.0 78.0
From Ref. [3].
(-LG), ␣-lactalbumin (␣-LA), bovine serum albumin (BSA), lactoferrin, lactoperoxidase, glycomacropeptide, etc., some of which could have important end uses, are getting disposed off as sewage. Even in WPCs and individual whey protein concentrates (IWPCs), the unique nutritional, therapeutic and functional characteristics of the individual whey proteins are largely unrealized due to interactions between components and degradation during processing. There, thus has been considerable commercial interest in the production of individual whey proteins with well characterized functional and biological properties by a process which will not denature, but retain its nutritional and other properties. Extensive researches have been carried out during last few years keeping this perspective in view [4,5]. -Lactoglobulin (-LG) is a better foam stabilizer than the other whey proteins and can be used in the production of confectionery, whereas ␣-lactalbumin (␣-LA) can be used in infant formula and as a neutraceutical because of its high tryptophan content. It also provides enhanced whippability in meringue like formulations. In addition, ␣-lactalbumin has strong affinity for glycosylated receptors on the surface of oocytes and spermatozoids and may, thus have strong potential as a contraceptive agent [6]. Extensive studies have been carried out by the researchers to demonstrate the feasibility of membrane systems for the separations of proteins with very similar molecular sizes. Van Reis et al. [7] used high performance tangential flow filtration (HPTFF) to separate bovine serum albumin (BSA) from an antigen binding fragment of a monoclonal antibody (Fab), achieving more than 900-fold purification and 90% yield of the BSA. In a previous work carried out by Zhang [8] regarding the study on the affinity cross-flow filtration process for separation of BSA from IgG molecules, the target protein BSA was recovered in more than 95% purity. van Eijndhoven et al. [9] developed a membrane system for the separation of bovine serum albumin (BSA) from hemoglobin, two proteins with essentially identical molecular weight, with more than 100-fold purification and nearly 70% yield. More recently, Cheang and Zydney [10] were able to obtain 100-fold purification and greater than 90% recovery of -LG from a binary mixture with ␣-LA. All these experimental data were obtained with model system consisting of an artificial mixture of two previously purified proteins. Experimental studies with complex multicomponent feed streams are much more limited, and the overall performance of these systems is much less impressive. With an objective to purify ␣-LA from acid casein whey, Muller et al. [11] used a combined ultrafiltration–diafiltration process and an ␣-LA purity of
only 50% was obtained in the final permeate. Bottomley [12] described a two-stage membrane process for the purification of ␣-LA from cheddar whey, but the final product still contained nearly 25% -LG. Lucas et al. [13] demonstrated the extraction of ␣-LA selectively from acid casein whey protein concentrate (WPC) at pH 7 by limiting -LG transmission by chemically modifying the inorganic membrane by a polyethyleneimine coating bearing positive charges. Some studies of proteins fractionation are based on the physico-chemical environment and the charge of the solutes. Mehra and Donnelly [14] showed the possibility of fractionating clarified or dialysed whey with cellulose membranes. The retentate was enriched with the high molecular weight proteins (immunoglobulin, lactoferrin, bovine serum albumin) and the permeate was enriched with the low molecular weight proteins (-lactoglobulin, ␣-lactalbumin). Ion exchange membrane chromatography (IEMC) [15,16] is a high resolution separation technique most suited for protein purification protocols. The separation is based on the reversible electrostatic interaction between a charged protein molecule and the oppositely charged chromatographic membrane (ion exchanger). Amino acids, for example, have both acidic and basic properties and these charge properties have a definite role on the reactivity of proteins to ion-exchange media. An ionexchanger consists of an insoluble matrix, such as cellulose, silica or often a form of polymer, styrene divinylbenzene to which charged groups have been covalently bound. The charged groups are associated with mobile counter-ions. The counterions may be reversibly exchanged with other ions of the same charge. The protein must displace the counter ions and therefore bind to the exchanger on the membrane. Before proceeding to choose the right IEMC column, one must consider whether the target protein would be bound as polyanion or polycation. As proteins are usually handled in buffer media, one is usually interested in their isoelectric points (IEP). The protein is positively charged below IEP and it will bind to any cationic exchange membrane. At a pH higher than the IEP of the protein, the target protein will be negatively charged and bind to anionic exchange membrane. Desorption of biomolecules from the ion-exchange membranes begins after increasing the ionic strength or changing the pH of the elution buffer. The work in this paper has been undertaken with an objective to examine the relative separation of -LG from whey protein concentrate using two-stage ultrafiltration with 30 and 10 kg mol−1 MWCO membranes in a stirred rotating disc membrane module, followed by ion-exchange membrane chromatography. It is reported [17] that at 20 ◦ C, bovine -LG forms a dimer at neutral pH, while the monomeric native state is stable at pH values below 3 and at low ionic strength in absence of salts. To study the overall performance at two different pH, one above and the other below the IEP of casein (IEP of casein being 4.6 [5]), experiments were carried at two different pH, i.e. 2.8 and 5.6. In each stage of ultrafiltration, the effects of stirrer speed, membrane disc rotation, transmembrane pressures (TMP) and solution pH on permeate flux and rejection were thoroughly investigated. The results obtained were compared in terms of the concentration ratio of -LG to ␣-LA and separation factor, in addition to other variables like flux, rejection, etc. One thing
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that needs to be mentioned is that, in IEMC spin column, the driving force is centrifugal which limits its use and scale-up in larger industrial version. But here, the main emphasis is to show the feasibility of the concept, i.e. possibility of obtaining a relative separation between ␣-LA and -LG using anion exchanger membrane, rather than to suggest a process flow-sheet which is directly applicable in industry. 2. Materials and methods 2.1. Materials Experiments were performed using casein whey obtained from local sweet meat industries situated in-and-around Kolkata, India. The pH of the raw casein whey was found to vary from 3 to 4, depending upon the quantity of excess acid present in whey resulting from acid caseination. In most of the cases, the sweet-meat industries use hydrochloric acid or its equivalent, for casein precipitation. Several standard proteins like ␣-lactalbumin (␣-LA), lactoglobulin (-LG), bovine serum albumin (BSA), lactoferrin (LF) and immunoglobulin (Ig) were purchased from either E.Merck (Mumbai, India) or from Sigma–Aldrich (St. Louis, USA) for forming standard solutions, for characterization of high performance liquid chromatography (HPLC) readings. All other chemicals (required for membrane cleaning, storage, and HPLC measurements) were purchased locally, mostly from E.Merck (Mumbai, India). The deionized water used in all the experiments was obtained from Arium 611DI ultrapure water system (make: Sartorius AG, G¨ottingen, Germany). The feed to
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this Arium 611DI was taken from usual laboratory distillation unit. 2.2. Membrane and module Ultrafiltration of pretreated casein whey was carried out batch wise in stirred rotating disc module (capable of being used as fixed disc module also) using polyether sulfone (PES) membrane. The module made of SS316, was manufactured by Gurpreet Engineering Works, Kanpur, UP (India) as per specified design. The module was equipped with two motors (Fig. 1) with speed-controllers to provide rotation of the stirrer and membrane housing. The module has the facility to rotate the membrane and the stirrer in opposite direction to provide maximum shear in the vicinity of the membrane. Digital tachometer was used to measure the rotational speed of both the membrane and the stirrer. Adequate mechanical sealing mechanism was provided to prevent leakage from the rotating membrane assembly. The magnetic drive stirrer mechanism prevents any possibility of leakage from the top stirrer. The complete schematic diagram of the rotating disk module setup was given in Fig. 1. The PES membranes (flat disk of 76 mm diameter) of 5, 10 and 30 kg mol−1 molecular weight cut-off (MWCO) were imported from Millipore Corporation, Bedford, MA (USA) through its Indian counterpart (Millipore India Limited). The 5 kg mol−1 membrane was used for the DD operation, which was also carried out in the same stirred rotating disk module, operated under a fixed set of operating conditions. The flat sheet membrane, operable in pH range of 1–14, has an actual diameter of 76 mm whereas the effective diameter was 56 mm. The membrane is
Fig. 1. Schematic diagram of rotating disc membrane module setup.
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Table 2 Preparation of loading buffer (of volume 100 mL) of different pH
2.3. Pretreatment of casein whey
Loading buffer pH
Amount of 0.1 M citric acid solution (mL)
Amount of 0.1 M trisodium citrate solution (mL)
4.6 4.8 5.0
44.5 40.0 35.0
55.5 60.0 65.0
In order to prevent any possibility of membrane fouling, the suspended casein particles and colloidal matters (mostly fat) were removed by centrifugation followed by microfiltration. Centrifugation was carried out in a research centrifuge model TC 4100D (make: Remi, Mumbai, India) with a speed of 12,500 rpm giving a RCF (relative centrifugal force) of 16,000 g for a period of 30 min. After centrifugation, the sample was subjected to microfiltration (MF) using “all glass vacuum filtration unit” (make: Sartorius AG, G¨ottingen, Germany), fitted with oilfree portable vacuum pump (Sartorius AG, G¨ottingen, Germany, model ROC 300 with moisture trap), with polyether sulfone (PES) membrane (47 mm diameter, pore size 0.45 m, Sartorius AG, G¨ottingen, Germany) being used as filter media. PES membranes exhibit no hydrophobic or hydrophilic interactions and are usually preferred for their low fouling characteristics, broad pH range and durability. The permeate from the microfiltration still contains lactose and different minerals, which needs to be removed to obtain, so called whey protein concentrate (WPC). The removal of lactose is very much essential as it may impair the downstream separation operation. Lactose and all other mineral salts were removed by a four-stage discontinuous diafiltration (DD) with 5 kg mol−1 PES membrane, the detail procedure of which will be mentioned later. The final permeate from the DD was adjusted for pH by adding calculated amount of hydrochloric acid or sodium hydroxide, as required to produce the feed for the subsequent UF run at two desired pHs, i.e 2.8 and 5.6.
compatible both in alcohol and aqueous medium and can be sterilized by immersion in 3% formalin or 5% hydrogen peroxide or 0.1% peracetic acid for 24 h. The retentate obtained after carrying UF using 10 kg mol−1 membrane (which mainly contains only -LG and ␣-LA) was subjected to ion-exchange membrane chromatography (IEMC) using VivapureTM Q Mini-H column (manufactured by Vivascience AG, Hannover, Germany; supplied by Sartorius Biotech Pvt. Limited, Bangalore, India) which has a strong basic anion exchanger with quaternary ammonium ion [R CH2 N+ (CH3 )3 Cl− ] immobilized on regenerated cellulose acetate substrate. This membrane was housed in a spin column with centrifugal action being used as the driving force. For equilibration of the membrane, 25 mM Tris/HCl, 25 mM NaCl buffer pH 8.0 was supplied with VivapureTM Q Mini-H column. Loading buffer of different pH was prepared by mixing citric acid solution and trisodium citrate solution. Citric acid monohydrate and trisodium citrate dihydrate were purchased from E-Merck, Germany. During experimentation, three different levels of loading buffer pH were selected to study optimum recovery of -LG. The levels were 4.6, 4.8 and 5.0, and the method of preparation of these loading buffer were mentioned in Table 2. Complete process flow diagram showing all the steps carried out during this analysis has been shown in Fig. 2.
2.4. Membrane compaction and water run Prior to use in experiments, each membrane was subjected to compaction for about 1 h with ultra-pure water at a pres-
Fig. 2. Schematic diagram of the proposed process flow diagram.
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sure of 9 × 105 Pa, higher than the highest operating pressure used in this study, to prevent any possibility of change of membrane hydraulic resistance during ultrafiltration. Once the water flux becomes steady with no further decrease, it was concluded that full compaction of the membrane has taken place. After compaction, membrane hydraulic resistance (Rm ) was determined based on water runs at different transmembrane pressures (TMPs) of (4–7) × 105 Pa. Membrane was rinsed thoroughly with distilled water after every run with casein whey so as to remove any deposited fouling layer, which was followed by water runs to determine the extent of fouling. The drops in water fluxes obtained from such studies were found to be within 2% of initial water flux, thus showing minimum fouling resulting from the proposed separation scheme. 2.5. Methodology The centrifuged and microfiltered whey was batch diafiltered four times using 5 kg mol−1 polyether sulphone (PES) membrane in the above module at 500 rpm stirrer speed, 300 rpm membrane rotation speed and TMP of 4 × 105 Pa with an initial volume 350 mL. During first stage of DD, the operation was carried out until a volume concentration ratio (VCR) of 2 was achieved. The required VCR was attained in approximately 1.2 h, after which 175 mL of deionised water was added and second stage diafiltration was carried out. This procedure was continued in all the four-stages of diafiltration operation. Ultrafiltration experiments were carried out batch wise in rotating disc membrane module at different operating conditions starting each time with a pretreated and diafiltered feed volume of 350 mL (175 mL of deionised water was added to final DD retentate to makeup the volume up to 350 mL). As per the objectives of the experiments, the solution pH was varied at 2.8 and 5.6, and the TMP was fixed at (4–7) × 105 Pa. The stirrer speed (ns ) was set as 0 (fixed stirrer), 500 and 1000 rpm, whereas the membrane rotation speed (nm ) was varied at 0 (fixed membrane), 300 and 600 rpm. The stirrer speed and membrane rotation speed were controlled through speed controllers, regulating the power supply to the respective motors. The stirrer speed and the membrane rotational speed were checked with the help of a digital tachometer working on stroboscopic principle during experiment. The experiments were designed in such a manner so as to get the nature of variation of flux and rejection with each of the independent variables as mentioned above. The permeate of 30 kg mol−1 membrane was again subjected to UF in rotating disc module using 10 kg mol−1 membrane and all the experimental runs were taken at different values of TMPs, stirrer speeds, membrane rotation speeds and pHs, as mentioned earlier. After each experiment, the module was disassembled first, the membrane disc was thoroughly rinsed for 20 min under running water in laboratory wash basin, soaked in 1 L of 0.5 mM NaOCl in 0.5 M NaOH for 30 min in a beaker, subsequently soaked in deionised water for 5 h and then again thoroughly washed for 20 min. In each case, the water flux was found to regain by more than 98% of its original value, suggesting the cause of flux decline to be either osmotic pressure limited or due to reversible fouling layer.
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The two types of retentate of 10 kg mol−1 membrane, resulting from initial feed solutions at two different pHs (2.8 and 5.6) were then fed to ion-exchange membrane chromatography unit. Ion-exchange membrane chromatography (IEMC) is a high resolution separation technique most suited for protein purification protocols. The separation is based on the reversible electrostatic interaction between a charged chromatographic membrane (ion exchanger) and the target protein molecule. In this work, VivapureTM Q Mini-H column manufactured by Vivascience AG (Hannover, Germany), supplied by Sartorius Biotech Pvt. Limited (Bangalore, India) was used for this purpose. The unit comprises of a strong basic anion exchanger, in the form of quaternary ammonium ion (R CH2 N+ (CH3 )3 Cl− ), with a membrane of bed volume 240 L, membrane surface area 7.48 × 10−4 m2 and amount of ionic groups to be 145–218 equiv. mL−1 for monovalent ions, and the membrane supported on stabilized regenerated cellulose acetate matrix. Following steps were involved in the proposed ion-exchange membrane chromatographic separation: (1) Equilibration of the VivapureTM Q Mini-H spin column with 400 L of 25 mM Tris/HCl and 25 mM NaCl buffer at pH 8.0, and subsequent spinning at 2000 g for 5 min in fixed angle centrifuge. (2) An amount of 250 L loading buffer of desired pH, prepared from citric acid monohydrate and trisodium citrate dihydrate (see Table 2) was added to 50 L of sample and vigorously mixed. (3) An amount of 300 L diluted sample was loaded onto VivapureTM Q Mini-H column and rotated at 2000 g for 5 min. Flow-through-fraction sample contained enriched LG. (4) An amount of 50 L of above mentioned buffer of desired pH was loaded again to ensure complete recovery of unbound -LG. (5) The bound protein ␣-LA was recovered by loading 300 L of 25 mM Tris/HCl, 1 M NaCl buffer pH 8.0. The pH of loading buffer was selected at a pH 4.6, 4.8 and 5.0 with an understanding that ␣-LA (pI 4.5–4.8), would be negatively charged at this pH and would bind to the membrane whereas -LG being positively charged (pI being 5.2–5.4) would pass through the membrane surface. The filtrate thus obtained from IEMC was analyzed further for individual protein concentration measurement. Another important factor in this analysis was to study the repeatability/reproducibility of the results obtained through the proposed separation train. For this, all the experiments were repeated under the same operating conditions and in the same modules at least thrice (in some cases, even six to seven times), to assess the reproducibility of the results. In all the cases, deviations between the various runs were found to be minimal, mostly remaining within ±5%. In some instance, greater deviations of one or two points were indeed obtained, which may be due to the experimental error. The repeatability of the results have been indicated in various figures in the form of error bars, where those one or two points showing greater than 10% deviation have been
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considered as off-shoot points and were not included in the calculation of error bars. In some of the figures, the error bars were not shown to preserve the clarity and usability of the figures. 2.6. Analysis The total solid concentration (TS) was determined gravimetrically after drying the samples in an incubator at 105 ◦ C and measuring the weight of the residue. Total protein concentration (TPC) was measured using a Hitachi dual beam UV–vis spectrophotometer (Model No. U-2800, with UV solution software) by the principle of Bradford protein assay at 590 nm. The assay is based on the observation that the maximum absorbance for an acidic solution of Coomassie Brilliant Blue G-250 shifts from 465 to 590 nm when binding to protein occurs [18]. Regarding the individual protein concentration, the focus in this study was mainly to measure ␣-LA and -LG. High performance liquid chromatography (HPLC) was used in this study to measure individual protein concentration. Water® (model 600 S) (imported from Waters Associates, Milford, MA, USA) HPLC with Shodex size exclusion column (8 mm × 300 mm, MW range: 100–150,000 Da) was used for the experimentation purpose. The HPLC was operated with phosphate buffer (pH 6.8) and the flow-rate was fixed at 1 mL min−1 . Disodium hydrogen phosphate (Na2 HPO4 ) and sodium dihydrogen phosphate (NaH2 PO4 ) solutions were prepared each of 0.1 M concentration and 0.15 M sodium chloride (NaCl) was added to these mixtures. These phosphate buffers were used in the ratio of 40/60 and good separation was observed. Detection was done at UV-210 nm. The elution times in most of the cases were within 15 min. In this study, individual protein concentrations were also determined using reversed phase-high performance liquid chromatography (RP-HPLC) according to a modification of the method suggested by Resmini et al. [19] with a silica based wide pore C18 column (15 m, 3.9 mm × 300 mm) in the same Water® Gradient HPLC. In this case, detection was done at UV280 nm and the flow rate fixed at 1 mL min−1 . The mobile phase was a binary system of water with 0.1% of trifluoro acetic acid (TFA) (buffer A) and acetonitrile with 0.1% of TFA (buffer B). The gradient condition was consisted of a two-step linear binary gradient: buffer A/buffer B = 70/30–35/65 (vol.%), gradient time of 15 min. The second method of determination produced a little bit better separation of the protein concerned (i.e. ␣-LA and -LG), though in this study, most of the results were obtained using the first method of size exclusion chromatography. In this study, during interpretation of IEMC results, a parameter called ‘separation factor’ (S) defined by: S=
(-LG : ␣-LA)filtrate (-LG : ␣-LA)feed
(1)
has been used, where, subscript filtrate and feed corresponds to those in ion-exchange chromatography membrane. Separation factor was plotted against loading buffer pH used in VivapureTM Q Mini-H column.
Fig. 3. Flux in four-stage discontinuous diafiltration operation.
3. Result and discussion After membrane compaction, series of water runs were taken to evaluate membrane hydraulic resistance Rm which was found to be (8.264 × 1012 ± 3.7 × 1011 ) m−1 for 30 kg mol−1 membrane and (1.998 × 1013 ± 9.6 × 1011 ) m−1 for 10 kg mol−1 membrane. The higher value of membrane hydraulic resistance for 10 kg mol−1 membrane compared to 30 kg mol−1 membrane is due to reduced pore size and compact nature of the membrane. The variation of flux with cumulative time during four-stage diafiltration with 5 kg mol−1 PES membrane was depicted in Fig. 3. It can be observed from Table 3 that only the pretreatment step alone, consisting of centrifugation, microfiltration and diafiltration, produced an improvement of %-purity of -LG with respect to total protein from 45.2% to 52.4%. The most important is the reduction of total solid (TS) in the pretreatment stage, from 645 kg m−3 in the raw casein whey to 13.89 kg m−3 in the pretreated casein whey, which accounts for 97.8% reduction in total solids. The major reduction in the total solid could be attributed to diafiltration step, which is very much suitable for purification of high molecular weight fraction. Due to reduction in total solid, the whey protein fraction on dry basis (with respect to total solid) has improved significantly, from 0.54% in the raw casein whey to 23.2% in pretreated casein whey. It may be observed from Fig. 3 that during the first stage of diafiltration, the flux has reduced from 140.6 to 11.2 L m−2 h−1 after about 1.2 h when a VCR of 2 was achieved. The protein fraction could even be increased much more by continuing the diafiltration for additional stages and using a little higher VCR. But in this study, that was not our objective, the emphasis was mainly to study the relative separation between ␣-LA and -LG using IEMC and the effect of feed pH on the whole separation train. Detailed results of the suggested two-stage UF employing 30 and 10 kg mol−1 membrane are shown in Table 3 under constant operating condition (pH 2.8, TMP = 4 × 105 Pa, ns = 0 rpm and nm = 300 rpm). Along with the UF results, this table also highlights the findings from IEMC. Table 3 gives the concentration of ␣-LA, -LG and TPC in pretreated casein whey, as
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Table 3 Compositions of pretreated casein whey, UF permeate (30 and 10 kg mol−1 ) and UF retentate (10 kg mol−1 ) at different pH, under fixed operating conditions (TMP = 4 × 105 Pa, ns = 500 rpm and nm = 0 rpm) and results of IEMC with UF-10 kg mol−1 retentate Item
Concentration in raw casein whey (kg m−3 )
Concentration in pretreated casein whey (kg m−3 )
pH
␣-LA
1.26 ± 0.0466
0.93 ± 0.032
-LG
3.46 ± 0.13
3.23 ± 0.12
TPC
7.66 ± 0.29
6.16 ± 0.23
TS
645 ± 25
2.8 5.6 2.8 5.6 2.8 5.6 2.8 5.6 2.8 5.6
13.89 ± 0.5
%-Purity of -LG 45.2 ± 2.1 52.4 ± 2.6 with respect to TPC %-Purity of -LG 0.54 ± 0.032 23.2 ± 1.4 with respect to TS a These concentrations refer to those in the diluted samples.
UF permeate (30 kg mol−1 )
2.8 5.6
well as those in UF-30 kg mol−1 permeate, UF-10 kg mol−1 permeate and UF-10 kg mol−1 retentate. In addition to the protein concentrations, Table 3 also presents total solid (TS) at different stages of filtration, along with %-purity of -LG on total protein, as well as on total solid basis. Table 3 also highlights the effect of feed pH on concentrations of different proteins, as well as on TS and TPC. It could be observed that all the permeate concentrations were greater at different stages at pH 2.8, compared to that at 5.6. It is already reported [17], that bovine -lactoglobulin assumes a monomeric native structure at pH 3 in the absence of salt. This fact is also supported by another article [20] wherein it is also mentioned that at pH 6.2 -LG exists as dimer. Hence greater rejection of -LG at pH 5.6 is attributed due to forward shift of monomer–dimer equilibrium, thereby causing formation of major proportion of -LG which is retained by the membranes (10 and 30 kg mol−1 ) resulting in lesser -LG concentration in UF permeate as well as reduced TS and TPC values at all subsequent stages. The effects of transmembrane pressure (TMP) on flux at pH 2.8 have been shown in Fig. 4 with stationary membrane and stirrer rotating at 500 rpm. Several models were tested to find which fitted the transient variation of flux with respect to time under these conditions of operations. Statistical analysis proved that the relationship between flux and time could be best described by the model given by Sheppard and Thomas [21] based on minimum value of standard error and maximum value of correlation coefficient: J = J1 t −b
Concentration (kg m−3 )
0.81 0.80 2.73 2.29 3.87 3.51 11.50 11.15 70.6 65.1
± ± ± ± ± ± ± ± ± ±
0.028 0.025 0.1 0.09 0.15 0.14 0.41 0.39 4.2 3.9
23.7 ± 1.64 20.5 ± 1.22
UF permeate (10 kg mol−1 ) 0.32 0.28 0.43 0.21 0.83 0.71 7.80 7.66 52. 4 30.4
± ± ± ± ± ± ± ± ± ±
0.011 0.009 0.013 0.007 0.03 0.024 0.31 0.31 3.1 1.9
5. 6 ± 0.38 2. 8 ± 0.18
UF retentate (10 kg mol−1 ) 1.02 0.99 3.71 3.05 5.11 4.55 15.20 14.64 72.7 67.0
± ± ± ± ± ± ± ± ± ±
0.035 0.034 0.14 0.11 0.19 0.17 0.61 0.54 4.5 4.6
24.4 ± 1.46 20.8 ± 1.45
IEMCa permeate (buffer pH 5.0) 0.031 0.028 0.539 0.437 0.615 0.502 2. 050 1.975 87.6 87.0
± ± ± ± ± ± ± ± ± ±
0.001 0.0008 0.02 0.017 0.024 0.018 0.077 0.098 5.25 5.1
26.3 ± 1.55 22.1 ± 1.32
sion parameters, J1 and b. Standard error was found to be 2.1417 and the correlation coefficient being 0.9979. Similarly, J1 and b was computed at other operating conditions. For all the runs, high values of correlation coefficients (>0.95) signifies good fitting of flux versus time data by Sheppard and Thomas model. The flux was observed to decline with time, which is obvious in any pressure-driven membrane process like UF; a 67% reduction of permeate flux was observed within a time span of 20 min at a TMP of 4 × 105 Pa whereas a reduction of 74.5% was observed within the same time period at a TMP of 7 × 105 Pa. This is probably due to the fouling caused by the whey proteins mainly BSA, IgG and partly ␣-LA and -LG onto the membrane surface because of possible formation of gel layer [2]. From the inset graph of Fig. 4, a 36% higher flux was obtained with a feed pH of 2.8 compared to the flux obtained with feed
(2)
where J is the volumetric flux in L m−2 h−1 measured at time t in seconds. The numerical values of J1 and b corresponding to the 4 × 105 Pa TMP and the specified values of nm and ns as shown in figure, have been estimated as 422.99 ± 11.42 and 0.3107 ± 0.0065, respectively, for the aforementioned units of flux and time. Nonlinear regression technique employing method of least-square was used to find the values of the regres-
Fig. 4. Effect of transmembrane pressures on whey flux at constant ns and constant nm using 30 kg mol−1 membrane. Effect of feed solution pH on flux vs. TMP plot at specified TMP, ns and nm .
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pH of 5.6 at a TMP of 5 × 105 Pa. This trend is valid at other operating conditions also. -LG exists in a dimmer–monomer equilibrium between pH 3.5 and 6.5 and as slight expanded monomer below pH 3.5 and above 7.5. Detail effects of pH on -LG and the resulting monomer–octomer–dimer equilibrium have been discussed elsewhere [17,22,23]. It has been already mentioned that, after four stages of discontinuous diafiltration of raw casein whey, the TS value in the pretreated casein whey has been reduced by 97.8%. Hence, the presence of minerals and salts may be considered to be almost negligible in the pretreated casein whey. Therefore, it may be stated that the effect of pH on protein charge was predominating at this low ionic strength. Apart from the presence of major fraction of -LG in monomeric form at pH 2.8, the less dense nature of deposited protein layer onto membrane surface at this pH is another reason for higher flux compared to that at pH 5.6. At pH 2.8, the magnitude of charge on -LG is more compared to that at 5.6 (IEP being 5.2–5.4). Due to stronger [24] mutual electrostatic repulsion among charged -LG molecules at pH 2.8, the deposited protein layer onto membrane surface is less dense. As a result, the concentration boundary layer is less dense causing the effect of polarized layer resistance to a lower value. Fig. 5 shows the variation of permeate flux as a function of time at different membrane rotation speeds under constant pH, ns and TMP. The decline of flux with time is obvious for any pressure-driven membrane processes like UF, which is due to concentration polarization and reversible or irreversible fouling. The flux enhancement with increase in membrane rotation speed could be observed from the figure, which resulted from reduced concentration polarization for a rotating membrane disc. It was observed that flux increased by nearly 33% at 1 min for a rotating membrane (300 rpm) compared to the corresponding
Fig. 5. Effect of membrane rotation on flux vs. time plot during UF of whey at a particular pH using 30 kg mol−1 membrane at constant TMP, ns . (Inset) Comparison of effect of stirrer speed on flux between 30 and 10 kg mol−1 membranes at a specified time at constant TMP and pH.
value for stationary membrane. The concentration polarization and the resulting fouling for stationary membrane could become very severe as the UF process goes on, which reduced the flux very drastically. The flux-decline between 30 s and 20 min of operation for stationary membrane was 67% whereas that for rotating membrane at 300 rpm, it was 46%. Effect of stirrer speed on flux at a particular time (5 min, J5 ) for both the 30 and 10 kg mol−1 membranes have been shown in the inset of Fig. 5, which clearly shows the flux enhancement with stirrer speed, provided all other operating variables remain constant. Increased stirrer speed creates more turbulence near the vicinity of the membranes which subsequently reduces the polarized layer resistance and accordingly increases the flux. Another important fact that can be noticed from the figure is the significantly higher flux of 30 kg mol−1 membrane in comparison to 10 kg mol−1 membrane under same condition of operations. This is expected as 10 kg mol−1 membranes are much more compact with less average pore size, thus giving reduced flux in comparison to 30 kg mol−1 membrane. Variations of observed rejection (R) with TMP for different solution pH under fixed nm and ns are shown in Fig. 6 for UF operations with 30 kg mol−1 membrane. The rejection was higher for pH 5.6, compared to the corresponding results at pH 2.8, the reasons of which have already been discussed earlier. Under a fixed set of operating conditions, the average rejections were found to increase with TMP, which is in somewhat contradiction with the observation in many cases, where apparent rejection has been found to decrease with TMP. Higher TMP would result in higher convective flow and accordingly more
Fig. 6. Effect of TMP on rejection during UF using 30 kg mol−1 membrane at two different pH at constant stirrer speed and with stationary membrane. (Inset) Effect of membrane rotation speed on rejection vs. TMP at constant stirrer speed.
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Fig. 7. Plot of value of the ratio -LG to ␣-LA at four-stage of processing, i.e. pretreatment, first stage UF (30 kg mol−1 ), second stage UF (10 kg mol−1 ) and final stage, i.e. ion-exchange-membrane process.
solute would be transported towards the membrane surface. This would results in higher thickness of the polarized layer as well as high membrane surface concentration. In this case, the increase of rejection with TMP may be explained as follows. Higher pressure has caused more compaction of the polarized layer; the layer as a result has acted as a secondary membrane, thus increasing the rejection due to steric hindrance. Effect of membrane rotation on rejection is shown in the inset of Fig. 6. Membrane rotation was found to decrease the observed rejection, which could be attributed due to reduced polarized layer thickness, thus reducing the secondary membrane effect and increasing solute transport through the membrane. Fig. 7 depicts the pattern of improvement in the ratio of (LG:␣-LA) at different stages and at two different pHs. The first stage corresponds to the pretreated casein whey, second stage corresponds to UF permeate from 30 kg mol−1 membrane and third stage the UF retentate from 10 kg mol−1 membrane with the fourth stage being the permeate (eluted stream) from IEMC. The average values of the ratio of -LG to ␣-LA at each stage were calculated from Table 3 and reported at two different feed solution pH, i.e.2.8 and 5.6. It may be observed that the value of the above mentioned ratio in the permeate of 30 kg mol−1 membrane at initial feed pH 2.8 was 3.37, whereas the value of this ratio was 2.85 corresponding to feed solution pH of 5.6. The retentate of 10 kg mol−1 membrane corresponding to the permeate of 30 kg mol−1 membrane of initial pH of 2.8 was found to give a higher value (3.64) of the ratio of -LG to ␣-LA compared to corresponding value (3.07) of 10 kg mol−1 retentate obtained from 30 kg mol−1 permeate of initial feed solution having pH of 5.6. The reasons of such findings have been discussed earlier. Finally, the stage 4 corresponds to IEMC, the results of which have been shown for loading buffer pH of 5.0 (for which the highest separation was obtained). The ratio (-LG:␣-LA) has improved substantially during the final stage in the elution phase (permeate) of ion-exchange membrane chromatographic separation process. The ratio corresponding to the initial feed solution pH 2.8 was calculated as 17.15 which was substantially higher
149
Fig. 8. Plot of separation factor vs. pH of loading buffer in ion-exchange membrane chromatography (IEMC).
than the magnitude of the ratio (9.64) corresponding to feed solution of pH 5.6. In IEMC, for the case reported in Fig. 7, the pH was adjusted in both cases to 5.0 using loading buffer, but the difference in the ratio (-LG:␣-LA) solely resulted from the difference in -LG content in UF-10 kg mol−1 retentate (which was used as feed to IEMC) corresponding to two different pH values. The variation of separation factor as a function of the loading phase pH in the ion-exchange affinity membrane separation process has been shown in Fig. 8. It may be observed from Fig. 8 that a separation factor of 4.71 was obtained against a loading buffer pH of 5.0 whereas only 1.42 was obtained when the sample was mixed with loading buffer of 4.6. This is possibly due to the fact that at loading buffer pH 5.0, the ␣-LA in sample remains as strongly negatively charged (IEP of ␣-LA being 4.2) and was more likely to be bound to the strong basic anion exchanger membrane whereas almost all the -LG passes through the membrane. In contrast, at loading buffer pH 4.8, the negative charge on ␣-LA is less in comparison to that at pH 5.0; hence a lesser fraction of ␣-LA got bound to the membrane. This is probably the reason why separation factor values decreased rapidly as the pH was decreased from 5.0 to 4.6. 4. Conclusion In the present work, detailed investigations on fractionation of casein whey by performing ultrafiltration in rotating disc membrane module was done with an objective to purify -LG in final product. Several pretreatments like centrifugation, microfiltration and diafiltration were used to remove suspended casein, lipids, lactose and other minerals. Diafiltration was found to be very important and using four-stage diafiltration with VCR 2 in each stage, a large improvement of -LG purity with respect to total solid was observed. In fact, -LG purity increased from 0.536% to 23.2% using these diafiltration steps. After pretreatment, two-stage UF, using 30 and 10 kg mol−1 membrane was used, followed by IEMC. Effects of various independent variables in UF, like feed solution pH, transmembrane pressure, stirrer speed, membrane rotation speed were investigated in detail
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and suitable separation scheme has been formulated. Higher stirrer speed and membrane rotation speed justifiably gave higher permeate flux though their rotational speeds were needed to be optimized. Higher flux and lower rejection were obtained in UF with the membranes, 30 and 10 kg mol−1 , at pH 2.8, compared to that at pH 5.6. This could be explained based on the effect of pH on possible dimerization of -LG between pH ranges from 3.5 to 6.5 and protein charge at this low ionic strength, which has also been reported elsewhere. It was observed through this study that to obtain highest separation of ␣-LA and -LG, it is recommended to keep the solution pH of pretreated whey in the range where -LG would remain the monomer state. This would result in highest separation and at pH 2.8, the ratio -LG/␣-LA was found to be 17.15 compared to 9.64 at pH 5.6. Further, selection of loading buffer pH in IEMC was found to be very important and in this case, a loading buffer pH of 5.0 was found to give best separation of ␣-LA and -LG. This study resulted in good relative separation of ␣-LA and -LG, and -LG purity of nearly 87% on total protein basis was obtained, though the purity was quite low (26%) on total solid (dry) basis. To improve this percentage, the lactose and other minerals, which still remained after four-stage diafiltration steps need to be removed. This requires some further study, but it could be accomplished by increasing the VCR in diafiltration step and continuing for more steps in diafiltration. Acknowledgement All the works in this study has been performed on equipment purchased under AICTE (TAPTEC) project (No. 8021/RID/NPROJ/TAP-6/2002-03) and are the part of the corresponding project work. Accordingly, the contribution of All India Council for Technical Education (AICTE), Government of India is gratefully acknowledged. References [1] K. Scott, Handbook of Industrial Membranes, Elsevier, Oxford, 1995. [2] M. Cheryan, Ultrafiltration and Microfiltration Handbook, Technomic Publications, Lancaster, 1998. [3] L. Pedersen, J. Mollerup, E. Hansen, A. Jungbauer, Whey proteins as a model system for chromatographic separation of proteins, J. Chromatogr. B 790 (2003) 161–173. [4] E.E.G. Rojas, J.S.D.R. Coimbra, Size-exclusion chromatography applied to the purification of whey proteins from the polymeric and saline phases of aqueous two-phase system, Process Biochem. 39 (2004) 1751– 1759.
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