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Biochemical Engineering Journal 48 (2009) 1–5

Contents lists available at ScienceDirect

Biochemical Engineering Journal

journa l homepage: www.e lsev ier .com/ locate /be j

ecovery of protein from brewer’s spent grain by ultrafiltration

e-Song Tang a, Gang-Ming Yin a, Yuan-Zhe He a, Song-Qing Hu a, Bing Li a, Lin Li a,∗,ui-Ling Liang b, Devajit Borthakur c

College of Light Industry and Food Sciences, South China University of Technology, Guangzhou 510641, ChinaCollege of Horticulture, South China Agriculture University, Guangzhou 510642, ChinaDepartment of Biotechnology, Plant Improvement Division, Tea Research Association, Jorhat, Assam 785008, India

r t i c l e i n f o

rticle history:eceived 28 December 2008eceived in revised form 23 April 2009

a b s t r a c t

Application of ultrafiltration in recovery of protein from brewer’s spent grain (BSG) was studied in thiswork. The effectiveness in removing water and salts was evaluated. Results indicated that increasing ofcross-flow rate could improve the limiting flux. More than 92% of the protein was retained by the mem-

ccepted 27 May 2009

eywords:rewers’ spent grainroteinembraneltrafiltration

branes with both MWCO of 5 and 30 kDa. The protein contents in the final product were 20.09 ± 1.40%and 15.98 ± 0.58%, respectively by 5 and 30 kDa membranes compared with that of 4.86 ± 0.61% concen-trated by rotary evaporation. It indicated that ultrafiltration had good ability in the removal of salts inthe extract solution and improved the quality of final products. The 5 kDa membrane had a little higherprotein retention capacity than that of 30 kDa.

© 2009 Published by Elsevier B.V.

. Introduction

Brewers’ spent grain (BSG) is the major by-product of the brew-ng industry, representing around 85% of the total by-productsenerated. BSG has high content of protein and protein contentore than 20% on dry weight basis is reported [1]. BSG is of

ow cost and high nutritive value. The ingestion of BSG, or itserived products, has health benefits. Incorporation of BSG into ratiets is beneficial to intestinal digestion, alleviating both constipa-ion and diarrhoea. Such effects were attributed to the content oflutamine-rich protein, and to the high content of non-cellulosicolysaccharides and smaller amounts of �-glucans [1].

For a long time, the main application of BSG has been limiteds animal feed along with utilization of BSG in increasing bricksorosity [2], removal of Cu(II) ions from aqueous solutions [3], ands brewing yeast carrier [4,5]. The incorporation of BSG into ready-o-eat snacks was also studied [6,7]. Due to the presence of manyeneficial components in BSG, separation of BSG into its individ-al components for both food and non-food applications is found

mportant. These researches included valorization of BSG to recoveraluable compounds such as �-tocopherol by supercritical fluidxtraction (SFE) technology coupled with pretreatment processes8], recovery of ferulic acid in BSG by sequentially extracting with

∗ Corresponding author. Tel.: +86 20 87112214; fax: +86 20 87113252.E-mail address: felinli@scut.edu.cn (L. Li).

369-703X/$ – see front matter © 2009 Published by Elsevier B.V.oi:10.1016/j.bej.2009.05.019

alkali of increasing strength [9], solublilization carbohydrates fromBSG by microwave radiation to 160 ◦C in the presence of 0.1 M HCl[10], extraction of ferulic and p-coumaric acids by alkaline hydrol-ysis of BSG [11], recovery of lignin from BSG [12] and production ofoligosaccharides [13].

For the separation of protein, alkaline extraction and proteinprecipitated by the addition of ammonia sulphate [14] or by acid-ification to pH 4.5 using 4N HCl [15] were commonly employed.Salts residue was removed by dialyzing against distilled water. Celuset al. [16] used alkaline (17%, w/v) extraction with 0.1 M NaOH at60 ◦C. After 60 min of extraction, samples were filtered (180 �m)and the proteins in the filtrate were precipitated by acidification topH 4.0 using 2.0 M citric acid. The precipitated protein was obtainedafter centrifugation at 10,000 × g for 10 min at 4 ◦C and was finallyfreeze-dried. Diptee et al. [17] extracted protein from BSG using0.6% Na2HPO4 solution, and ethanol was added to precipitate pro-tein. But it was difficult to obtain protein from BSG in our previousexperiment since it was easily denatured by temperature duringprocessing. And the salt remained in the extract solution couldnot be efficiently removed by dialysis. Membrane separation hasmany advantages such as absence of phase change of water andreduced energy consumption for the removal of water and small

molecular size compounds, such as salts. Since no heat was addedto the co-product stream, quality of protein in the co-productsshould not be compromised [18]. In recent years, ultrafiltration hasbeen used for the separation of a wide range of compounds [18,19]and membranes are used extensively throughout the production

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f biotechnology products [20]. Application of membrane processight overcome the problem met in BSG protein isolation.

However, the major problems of ultrafiltration process are con-entration polarization and fouling which reduce the permeate fluxar below the theoretical capacity and change membrane selectiv-ty. Both concentration polarization and fouling strongly depend onperation conditions and membrane characteristics, such as feedroperties, membrane molecular weight cut off (MWCO), trans-embrane pressure (TMP) and cross-flow rate. In the present work,

pplication of ultrafiltration in removing water and salts in theecovery of protein from BSG was investigated and the effectivenessas evaluated.

. Materials and methods

.1. Preparation of protein solution

BSG (73.8% moisture, 7.6% protein, Kjeldahl N × 6.25, wet weightasis) was obtained from Guangzhou Zhujiang Brewery Group Co.,td., China. It was kept in −20 ◦C and thawed at 4 ◦C overnightefore extraction. The BSG extract was prepared by ultrasound-ssisted extraction using sodium carbonate buffer (pH 10), whichas found good in extraction of protein from BSG in our lab (unpub-

ished data). The protein solution was prepared by extraction of BSG100 g) using 1000 ml of extractant for 1 h and then filtered throughylon cloth and the filtrate solution was centrifuged (10,000 × g) at◦C. The supernatant was collected and used as feed solution.

.2. Equipment and membranes

Minipore LabscaleTM TFF system (Millipore, USA) was used forhis study. The system consisted of a re-circulation pump, cross-ow ultrafiltration module (Pellon-XL Module, Millipore, USA)quipped with membrane of BIOMAX®. The trans-membrane pres-ure and cross-flow velocity were adjusted by a manual valve andump controller. The pressure was measured by a standard pres-ure gauge. Membranes of MWCO of 5 and 30 kDa with a surfacerea of 0.05 m2 were used in these experiments.

All experiments were operated at ambient temperature∼25 ◦C). BSG extract (500 ml) was used as the feed for each exper-ment. The pressures were regulated using pressure gauges. Theross-flow rate and permeate solutions were measured using grad-ated cylinder and stopwatch.

After each run, the membrane was cleaned by alkali treatments recommended: a solution of 0.1 M sodium hydroxide was recy-led past the membrane at a cross-flow rate of 1.0 L/h for an hourith TMP of 25 psi. The storage solution recommended for this type

f membrane is 0.1 M sodium hydroxide and was used in all thexperiments.

.3. Effects of membrane MWCO and operating conditions

The effects of membrane MWCO and operating conditions weretudied using the total recycle mode. Both retentate and perme-te were re-circulated to feed tank. The trans-membrane pressureTMP) of 10–45 psi was used for both membranes. The cross-flowates were controlled at 0.9, 1.8 and 2.7 L/h for both membranes.he permeate flux was measured using a graduated cylinder and atop watch. The samples of permeate and feed bulk were collectedor protein analysis.

.4. Solution concentration

A single batch concentration was investigated for the removal ofolvent. The retentate was recycled to the feed bulk while the per-eate was removed from module. At each experiment, 0.5 L sample

eering Journal 48 (2009) 1–5

solution was concentrated. TMP was controlled at 25 psi. The cross-flow rate was controlled similar to those used in the total recyclemode. At each condition, the permeate flux was measured until per-meate flux was constant. The permeate and retentate samples werecollected for analysis. The concentrated solution was lyophilized toget final products.

A control concentration method of vacuum rotary evaporationwas carried out at 40 ◦C using a rotary evaporator (RE-52CS/5299,Shanghai Yarong Ltd., China). The protein solution with volumeof 0.5 L prepared from BSG was concentrated to 0.1 L, and thenlyophilized to get final protein products. The experiment repeatedin triplicate.

2.5. Analytical methods

Soluble protein concentration in the permeate and feed bulk wasmeasured by Bradford method [21], using bovine serum albumin(BSA) as the standard. One milliliter of diluted sample was placedin a test-tube. Five milliliters of Bradford dye was added and mixedand allowed to stand for 1–2 h. The absorbance of the mixed samplewas measured at 595 nm with a UV–vis spectrophotometer. Theconcentration of protein in the sample was determined using thestandard curve of UV absorbance and concentration.

The average pressure experienced by the membrane surfacebetween the feed and retentate ports is called the trans-membranepressure (TMP) and is calculated using Eq. (1):

TMP = Pin + Pout

2(1)

where Pin is the feed pressure (psi) and Pout is the retentate pressure.Membrane flux is a measure of the permeate flux taking into

account the active surface area of the membrane and is calculatedusing Eq. (2):

J = 1A

dV

dt(2)

where J is the permeate flux (L/m2 h), A is the area of the membrane(m2), V is the filtrate volume (L) and t is the unit time.

Total membrane resistance (Rm) can be determined from Eq. (3):

Rm = �P

�J(3)

where �P is the filtration pressure, it is equivalent to the TMP here;and � the solution viscosity, was determined by a viscometer (DV-IViscometer, Brookfield Engineering Labs, Inc., USA). The changes ofmembrane resistance for pure water were determined before andafter filtration.

The protein retention ratio (R) was defined as

R = 1 − CP

CF(4)

where CF is the concentration of protein in feed stream and CP isthe concentration of protein in permeate.

The yield was calculated as

Y(%) = Pfinal

PBSG× 100 (5)

where Pfinal is the protein content in final product and PBSG is theprotein content in BSG (wet basis).

3. Results and discussion

3.1. Pure water flux

Pure water flux was measured at the beginning of the exper-iment. The flux increased linearly with TMP within the testedpressure range, 5–50 psi (Fig. 1), as same as reported by Chollangi

D.-S. Tang et al. / Biochemical Engineering Journal 48 (2009) 1–5 3

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ig. 1. Water permeate flux against trans-membrane pressure for 5 and 30 kDaembranes (Pellon-XL Module Millipore, USA) at ambulant temperature (∼25 ◦C)

nd different cross-flow rates.

nd Hossain [22] and Casa et al. [23]. The flow rates were 0.9, 1.8nd 2.7 L/h of 5 kDa MWCO membrane and 1.8 and 2.7 L/h for 30 kDaWCO membrane, respectively. Fig. 1 also indicates that the flow

ate had little effect on the permeate flux for pure water.

.2. Effects of membrane MWCO, cross-flow rate andrans-membrane pressure

Effects of membrane MWCO, cross-flow rate and TMP weretudied using total recycle ultrafiltration mode. The retentate andermeate were re-circulated to feed tank. The permeate flux waseasured by a graduated cylinder and a stop watch.

Fig. 2 shows the steady state permeate flux versus TMP at dif-erent cross-flow rate. It is evident that an increase of cross-flowate caused a higher permeate flux. The permeate flux was TMP-ependent when TMP was lower and it became TMP-independent

f TMP was higher enough.The increasing of cross-flow rate also resulted in an increase of

he permeate flux for the membrane with MWCO of both 5 and

0 kDa. In this study, the limiting fluxes at cross-flow rate 0.9, 1.8nd 2.7 L/h for the 30 kDa membrane were 55.86 ± 2.63, 87.04 ± 2.7nd 107.21 ± 4.01 L/m2 h, respectively. And the limiting fluxes atross-flow rate 0.9 and 1.8 L/h for 5 kDa membrane were 46.4 ± 2.77

ig. 2. Steady state permeate flux versus trans-membrane pressure at differentross-flow rate.

Fig. 3. Permeate flux versus permeate volume during concentration at differentcross-flow rate.

and 57.92 ± 2.34 L/m2 h, respectively. At the same cross-flow rate,the membrane with MWCO of 30 kDa had a higher permeate fluxthan that with MWCO of 5 kDa.

The increasing of cross-flow rate resulted in a reduction in con-centration polarization and led to an increase in the permeate flux.The experiments showed that the permeate flux could be improvedfurther by increasing the cross-flow rate. Since filtration resistancecaused by concentration polarization and reversible fouling layercould be significantly decreased with increasing cross-flow rate[19]. It was interesting that no limiting flux for 5 kDa MWCO mem-brane was obtained under the experimental conditions when thecross-flow rate was 2.7 L/h. The reason was probably that com-pounds in the extract solution could not deposit on the surface of5 kDa MWCO membrane and the reversible fouling layer could noteasily formed under the experimental condition.

3.3. Permeate flux profile during concentration process

Concentration procedure was carried out at TMP of 25 psi andthe permeate flux profiles observed for both 5 and 30 kDa mem-branes at different cross-flow rate (Fig. 3). The permeate fluxdecreased with the increasing of permeate volume. And at thebeginning of the filtration process the flux decreased sharply, andlater the flux decreased slightly. These permeate flux profiles weretypical for ultrafiltration process. The concentration polarizationwas the major factor for rapid decrease of fluxes during the initialseveral minutes. The second stage, i.e. slight decrease of the flux,was due to protein adsorption and particle deposition on the sur-face of membrane or interior wall of the membrane pore. Severalfactors, viz. particle deposition, consolidation of fouling materialson the interior space of membrane or formation of cake layer causedthe long period of constant flux [19].

Fig. 3 also indicates that increasing of cross-flow rate resultedin improving of permeate flux. Permeate flux difference betweenthe two types of membrane rose with the increasing of cross-flowrate (Table 1). Increasing of cross-flow rate could improve permeateflux to a greater extent for 30 kDa membrane than that of 5 kDamembrane. It indicated that increasing of cross-flow rate properlycould improve the performance of membrane.

3.4. Retention of protein

Protein retention was studied in total cycle model. Table 2 showsthe effect of trans-membrane pressure and cross-flow velocity on

4 D.-S. Tang et al. / Biochemical Engineering Journal 48 (2009) 1–5

Table 1Permeate flux difference between 5 and 30 kDa membranes at three cross-flow rate during concentration procedure.

Permeate volume (ml) Permeate flux differencea between 5 and 30 kDa membrane (L/m2 h)

0.9 L/h 1.8 L/h 2.7 L/h

0 4.14 27.52 36.2250 0.46 24.39 35.59

100 0.98 25.19 32.73150 0.85 27.8 37.05200 2.44 27.23 35.28250 1.43 24.66 33.63300 0.53 24.29 29.74350 1.65 21.52 29.38

a Permeate flux difference = mean permeate flux of 30 kDa membrane − mean permeate flux of 5 kDa membrane.

Table 2Effect of trans-membrane pressure, cross-flow rate and MWCO on the retention of crude protein in total cycle model.

TMP (psi) Protein retention ratio (mean ± SD%)

5 kDa 30 kDa

0.9 L/h 1.8 L/h 2.7 L/h 0.9 L/h 1.8 L/h 2.7 L/h

10 96.88 ± 0.01 97.74 ± 0.03 – 94.37 ± 0.2 96.05 ± 0.18 96.57 ± 0.0615 95.78 ± 0.14 97.62 ± 0.05 96.95 ± 0.23 94.1 ± 0.62 96.18 ± 0.02 96.29 ± 0.1620 94.18 ± 0.07 96.76 ± 0.03 96.50 ± 0.04 94.01 ± 0.08 95.95 ± 0.12 95.72 ± 0.0225 94.56 ± 0.51 96.19 ± 0.18 96.13 ± 0.11 93.25 ± 0.17 95.73 ± 0.02 94.98 ± 0.073 98 ± 03 44 ± 04 52 ± 04 30 ± 0

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0 94.52 ± 0.00 96.12 ± 0.07 95.5 94.37 ± 0.03 95.62 ± 0.02 96.0 94.31 ± 0.12 95.56 ± 0.02 96.5 93.74 ± 0.26 95.37 ± 0.02 96.

he protein retention ratio. It can be seen that most of the protein isetained by the ultrafiltration membranes with both MWCO of 5 and0 kDa. Protein retention ratios for both membranes with MWCOf 5 and 30 kDa decreased slightly with the increasing of TMP. Theetention of protein for 30 kDa membrane was lower than that ofkDa membrane.

During concentration procedure, retention of protein ratiohanged little for 5 kDa membrane with the permeate volume ris-ng but it increased slowly in case of 30 kDa membrane (Fig. 4). It

as also evident that the of membrane resistance increased greatlyefore and after filtration (Fig. 5) and the membrane resistancef 5 kDa membrane was higher than that of 30 kDa membrane.

fter filtration the membrane resistance increased almost six times

or 30 kDa MWCO membrane while about three times for 5 kDaWCO membrane. Membrane fouling was led to the increasing ofembrane resistance and it was easier occurred for 30 kDa MWCO

ig. 4. Protein retention of 5 and 30 kDa membranes during concentration proce-ure.

.18 92.58 ± 0.39 93.57 ± 0.63 93.30 ± 0.66

.13 92.38 ± 0.08 94.08 ± 0.10 94.58 ± 0.03

.04 93.65 ± 0.03 94.50 ± 0.38 94.65 ± 0.11

.18 – 92.65 ± 0.78 94.78 ± 0.13

membrane in this study. It could be inferred that pore blockageor pore constriction was the main fouling mechanism. The possi-bility was that small fraction of the extract was of molecular sizebetween 5 and 30 kDa and the pore of 30 kDa MWCO membranecould be easily fouled than that of 5 kDa membrane.

3.5. Comparison with rotary evaporator concentration

In this work, protein solution concentration by ultrafiltrationand rotary evaporation was compared. The protein yield and pro-tein content were showed in Table 3. It was found that the proteinyield was not significantly different. But the protein content in

the final product varied greatly: the protein in the final prod-uct was 4.86 ± 0.61%, 15.98 ± 0.58% and 20.09 ± 1.40% (mean ± SD,n = 3) respectively for concentration by rotary evaporation, 30 and5 kDa membranes respectively. It indicated that many small molec-

Fig. 5. Total membrane resistance for pure water before and after filtration at TMPof 25 psi.

D.-S. Tang et al. / Biochemical Engine

Table 3Comparison of protein yield and protein content in final products.

Treatment Protein yield (%) Protein content (%)

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Biochem. 72 (1976) 248–254.[22] A. Chollangi, M.M. Hossain, Separation of proteins and lactose from dairy

otate evaporation 12.69 ± 1.59 4.86 ± 0.61oncentration by 30 kDa membrane 10.01 ± 0.36 15.98 ± 0.58oncentration by 5 kDa membrane 14.59 ± 1.02 20.09 ± 1.40

lar size compounds were removed by membrane concentration.n this study, the protein in BSG was extracted using sodium car-onate buffer. Sodium carbonate remained in the product preparedy rotary evaporation, and it could be removed by membrane sep-ration. Table 3 illustrates that the yield and purity of protein innal product concentrated by 5 kDa membrane were higher thatf 30 kDa membrane. It indicated that 5 kDa membrane had betterbility in retention of protein and removal of salts.

. Conclusions

The results from this study suggested that membrane had manydvantages in respect of the concentration procedure for proteinxtraction from BSG. More than 92% protein was retained by theltrafiltration membrane with MWCO of 5, 30 and 5 kDa mem-rane had a little higher protein retention capacity than that of0 kDa. Increasing of trans-membrane pressure and cross-flow rateaused higher permeate flux. Ultrafiltration had good ability in theemoval of salts in the protein extracted by salt buffer solution andmproving the protein content.

cknowledgments

This work was supported by China Postdoctoral Science Foun-ation (20080440757); Post Doctor Innovative Foundation of Southhina University of Technology; Industry-Study-Research Basechievement translating program, Department of Education ofuangdong Province (cgzhzd0704); Guangdong Province nature

cience fund research group program (05200617); Guangdongrovince brainstorm program (2007B020801001).

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