Separation and Purification Technology 89 (2012) 125–134

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Separation and Purification Technology

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Clarifications of stevia extract using cross flow ultrafiltration and concentrationby nanofiltration

Chhaya a, Sourav Mondal b, G.C. Majumdar a, Sirshendu De b,⇑a Department of Agriculture and Food Engineering, Indian Institute of Technology, Kharagpur 721302, Indiab Department of Chemical Engineering, Indian Institute of Technology, Kharagpur 721302, India

a r t i c l e i n f o a b s t r a c t

Article history:Received 9 December 2011Received in revised form 8 January 2012Accepted 9 January 2012Available online 16 January 2012

Keywords:SteviosideClarificationUltrafiltrationPermeate fluxNanofiltration

1383-5866/$ - see front matter � 2012 Elsevier B.V. Adoi:10.1016/j.seppur.2012.01.016

⇑ Corresponding author. Tel.: +91 3222 283926; faxE-mail address: sde@che.iitkgp.ernet.in (S. De).

Cross flow ultrafiltration was employed to clarify the pretreated stevia extract. Two practical modes ofcross flow ultrafiltration, namely, steady state under total recycle mode and batch concentration mode,were used. A detailed investigation of effects of the operating conditions on the permeate flux and per-meate quality was undertaken. It was observed that the significant flux enhancement was achieved withtransmembrane pressure drop and cross flow rate. Maximum 200% flux enhancement with cross flowrate and 140% with transmembrane pressure drop were attained in the range of operating conditionsstudied herein. Effects of cross flow rate on the permeate properties were marginal but that of the trans-membrane pressure drop was significant. Recovery of stevioside in the permeate was in the range of 30–56% for various transmembrane pressure drop and it was maximum for lower operating pressure,276 kPa. However, the recovery of stevioside decreased to 38% at 276 kPa pressure after 10 h of operation.Nanofiltration was employed to concentrate the ultrafiltered liquor. During nanofiltration, the ultrafil-tered feed was concentrated maximum twice at 1241 kPa and 1500 rpm of stirrer speed within 1 h ofoperation. Maximum overall (ultrafiltration followed by nanofiltration) purity and recovery of 60% isobtained for a particular set of operating conditions.

� 2012 Elsevier B.V. All rights reserved.

1. Introduction

Stevia is an herbaceous plant and it is mainly found in SouthAmerica. Chemicals belonging to glycoside family, for example, ste-vioside, rebaudioside A, B, D, E, dulcoside A and B are responsible forsweet taste of the leaves of this plant [1]. Out of these, stevioside oc-curs in maximum amount in stevia leaves and other sweetenerscomprise about 4–20% of the dried leaves [2]. Stevioside is 300times sweeter than sugar, thermally stable at high temperature(about 100–120 0C), non-calorific and safer for diabetic patients[3,4]. Thus, the use of stevioside as a natural sweetener in place ofsugar is having high demand. Therefore, extraction of this naturalsweetener from stevia leaves is an important technical challenge.

Traditional processes for extraction of stevioside include addi-tion of chelating agents, like calcium hydroxide, followed by crys-tallization [5]; centrifugation followed by treatment with calciumhydroxide and further treatment by ion exchange resin [6,7]; sol-vent extraction followed by ion exchange [8]; adsorption with zeo-lite CaX in fixed bed columns [9,10]. These processes are timeconsuming, expensive and sometimes are not suitable for ediblepurposes when the solvent like methanol is used for extraction.An additional challenge is removal of the added chemicals and

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traces of organic solvent. Membrane based technologies can be aviable alternative in this regard. Various advantages of membranebased processes include, no phase change; addition of no chemi-cals and ease of scaling up. Membrane based processes have al-ready been used for processing of various juice and beverages.Some of these are, apple juice [11], mosambi juice [12], pomegran-ate juice [13], carrot juice [14], pineapple juice [15], kiwi fruit juice[16], water melon juice [17], green coconut water [18], black tea[19], selective extraction of (�)epigallocatechin gallate from greentea leaves [20], etc.

Few uses of membrane based technologies for extraction of ste-vioside from stevia leaves have also been attempted. In most of thecases, hybrid membrane separation was reported. Fuh and Chiang,carried out extraction of stevioside using two routes [21]. Thesewere, (i) precipitation by inorganic salts; (ii) clarification by ultrafil-tration in multi-effect diafiltration mode followed by concentrationby reverse osmosis. 25,000 and 100,000 Da molecular weight cut offmembrane were used for clarification at 12 and 8.5 bar transmem-brane pressure, respectively and at 25 l/min flow rate. Reverseosmosis was carried out at 45 bar and 25 l/min flow rate and thesolution was concentrated 10 folds. Stevioside with higher puritywas obtained in membrane based processes. Zhang et al. adopteda membrane based process for processing of stevia extract [22].Microfiltration (using 0.35 m ceramic membrane at 104 kPatransmembrane pressure) was used as pre-treatment followed by

126 Chhaya et al. / Separation and Purification Technology 89 (2012) 125–134

ultrafiltration (using 2500 Da cut off membrane at 440 kPa pres-sure) for subsequent clarification. Finally, nanofiltration was usedfor further purification and concentration of stevia extract at510 kPa. However, they have not reported any data regarding per-meate flux and recovery or final concentration of nanofiltration.Modified Zeolite (NaX) was used for pre-treatment of crude steviaextract followed by microfiltration [23]. Approximately 80–100%clarification of stevia extract was achieved with microfiltrationmembranes of varying pore sizes and at different operating pres-sures. A comparison based on the use of commercial and taylor-made PES membrane for stevioside purification has been reported[24]. Using microfiltration, ultrafiltration and nanofiltration, 37%purity and 30% yield were achieved. However, the above studieslack the analysis of the effect of operating conditions on systemperformance.

Decline of permeate flux during the course of filtration is a ma-jor problem [25,26]. Membrane fouling is responsible for this. Twotypes of fouling normally occur, namely, reversible and irrevers-ible. Reversible fouling is due to a phenomenon, known as concen-tration polarization, i.e. accumulation of solutes over themembrane surface resulting to additional resistance against thesolvent flux. This fouling can be removed by adopting a suitablewashing protocol. On the other hand, during filtration, solutemay adsorb on the pore mouth or inside the pores, thereby, block-ing the pores partially and/or completely. This fouling cannot beremoved completely by washing the membrane and a fraction ofmembrane permeability is lost permanently. This is known as irre-versible fouling. It is well-established that the fouling phenomenainvolved and the consequent membrane flux decline are insepara-bly associated with the operating conditions. Therefore, by select-ing a suitable set of operating conditions and the mode ofoperation, one can reduce membrane fouling. Hence, the role ofoperating conditions is extremely important in membrane separa-tion processes. Cross flow mode of operation is quite effective inarresting the growth of polarized layer of solutes over the mem-brane surface due to the shearing action of the retentate flow.

The works related to application of membrane processes fortreatment of stevia extract as described earlier do not provideany details of the effects of the operating conditions on the perme-ate flux decline as well as the permeate quality during ultrafiltra-tion/nanofiltration/reverse osmosis of stevia extract. Theseaspects are extremely important for an efficient design of indus-trial-scale processing unit. The present study is therefore, takenup to bridge this gap. In this work, cross flow ultrafiltration wasconducted in two modes, namely, total recycle and batch concen-tration mode with a detailed investigation of the effects of operat-ing conditions. Ultrafiltered solution was concentrated usingnanofiltration in a stirred cell and the effects of transmembranepressure drop and the stirrer speed were investigated in detail.

2. Materials and methods

2.1. Materials

Dry stevia leaves powder was obtained from M/s, RAS AgroAssociates, Maharashtra, India. Distilled water was used as solventfor extraction process. M/s, Merck India Limited, Mumbai, India,supplied high performance liquid chromatography (HPLC) gradeacetonitrile and water. Standard stevioside of 98% purity was ob-tained from M/s, Sigma–Aldrich, USA.

2.2. Membranes

For ultrafiltration, a 30 kDa membrane of polyethersulfone withpermeability 4.4 � 10�11 m/Pa s has been used. This membrane

was selected after investigating the performance of 5, 10, 30 and100 kDa membranes for extraction of stevia. 30 kDa membranewas found to have the highest permeate flux and maximum recov-ery of stevioside in the permeate [27]. These membranes were sup-plied by M/s, Permionics Membranes Pvt. Ltd., Gorwa, Vadodara,India. Nanofiltration was conducted using 400 molecular weightcut off membrane consisting of a polyamide skin over a polysul-phone support is supplied by M/s, Genesis Membrane SepratechPvt. Ltd., Mumbai, India. The permeability of the nanofiltrationmembrane was 1.44 � 10�11 m/Pa s.

2.3. Experimental set up

Two different experimental set ups have been used. A cross flowultrafiltration set up for clarification of stevia extract using ultrafil-tration. A stirred experimental set up was used for conduction ofnanofiltration experiments.

2.3.1. Cross flow ultrafiltration set upA rectangular cross-flow cell, made of stainless steel, was de-

signed and fabricated. Two neoprene rubber gaskets were placedover the membrane forming the flow channel. The channel heightafter tightening the two flanges was found to be 3.0 � 10�3 m. Theeffective dimension of the membrane was 14.5 cm in length and5.5 cm in width. The cell consisted of two rectangular matchingflanges. The inner surface of the top flange was mirror polished.The bottom flange was grooved, forming the channels for the per-meate flow. A porous stainless steel plate was placed on the lowerflange that provides mechanical support to the membrane. Twoflanges were tightened to create a leak proof channel for conduct-ing experiments in cross flow mode.

The centrifuged extract was pumped by a high pressure recipro-cating pump from the stainless steel feed tank to the cross flowcell. The retentate stream was recycled to the feed tank routedthrough a rotameter. The pressure and the cross flow rate insidethe membrane channel were independently set by operating thevalves in the bypass line and that at the outlet of the membranecell. Permeate samples were collected from the bottom of the celland were analyzed for color, clarity, total solids and stevioside con-centration. The membrane module assembly is presented in Fig. 1.

2.3.2. Stirred batch cell (for nanofiltration)The experimental set-up consisted of a stirred batch cell of

650 ml capacity made of stainless steel and it was pressurisedusing a nitrogen cylinder. For a typical run, about 300 ml of ultra-filtered feed was charged into the batch cell. The stirrer speed wasset using a variac (variable AC transformer for smooth control ofvoltage and thereby controlling the stirring speed) and was mea-sured by a hand held digital tachometer (Agronic, India). Insidethe cell, a circular membrane was placed over a base support.The membrane diameter was 6.6 cm and the effective membranearea was 34.2 cm2. The permeating solution from the bottom ofthe cell was used for further analysis. The schematic diagram ofthe experimental set-up is shown in Fig. 2.

2.4. Methods

2.4.1. Extraction processDry stevia leaves powder was mixed with hot distilled water at

a ratio of 1:14(g:ml). Temperature was fixed at 78 ± 1 �C for56 min. The above operating conditions were selected based onthe optimization experiments carried out for maximum extractionof stevioside in the liquor [28]. Constant temperature water bathwas used for hot water extraction. Next, the aqueous stevia extractwas cooled to room temperature and cloth filtered. The filtered

Fig. 1. Schematic diagram of the cross flow set up: (a) two flanges; (b) the top view of grooved bottom flange; (c) top view of bottom flange after putting the membrane.

2

3

1

4

1: Filtration cell. 2: Electronic balance. 3: Beaker for collection of permeate 4: Nitrogen cylinder

Fig. 2. Schematic diagram of the stirred batch cell.

Chhaya et al. / Separation and Purification Technology 89 (2012) 125–134 127

extract was analyzed for their color, clarity, stevioside concentra-tion and total solid content.

2.4.2. Primary clarificationPrimary clarification of the extract was carried out in a labora-

tory centrifuge (Model number R-24, supplied by M/s, Remi Inter-national Ltd., Mumbai, India). The centrifugation capacity was200 ml per batch. The operating conditions were stirrer speed5334 g and centrifugation time 26 min. These operating conditions

were selected based on the related optimization study so that max-imum clarity and recovery of stevioside were obtained.

2.4.3. Cross flow ultrafiltrationA fresh membrane was compacted at a pressure higher than the

maximum operating pressure for 3 h using distilled water and thenits permeability was measured. The extract was placed in a stain-less steel feed tank of 3 l capacity. A high pressure reciprocatingpump was used to feed the effluent into the cross-flow membranecell. Cumulative volumes of permeate were collected during theexperiment. Permeate samples were collected at different timeintervals for analysis. A bypass line was provided from the pumpdelivery to the feed tank. Retentate and bypass control valves wereused to vary the pressure and flow rate accordingly. Values of per-meate flux were determined from the slopes of cumulative volumeversus time plot. The precision of flux measurement was in the or-der of ±5%. The permeate stream after collecting required amountof sample was recycled to the feed tank to maintain a constant con-centration in the feed tank under total recycle mode. The permeatewas not recycled under batch concentration mode of operation.Duration of the cross-flow experiments was 45 min for total re-cycle mode of operation and it was 10 h for batch concentrationmode. The feed volumes were 2 l for total recycle mode and the ini-tial feed volume was 1.8 l for batch concentration mode. The sche-matic of cross flow experimental set-up is given in Fig. 3.

Once an experimental run was over, the membrane was thor-oughly washed, in situ, with distilled water for 30 min applying a

Fig. 3. Schematic diagram of cross flow set up.

128 Chhaya et al. / Separation and Purification Technology 89 (2012) 125–134

maximum pressure of 200 kPa. The cell was dismantled and themembrane was rinsed with distilled water and was dipped in 2%sodium dodecyl sulphate solution overnight. Next, the membranewas washed carefully with distilled water to remove traces of sur-factant. The cell was reassembled and the membrane permeabilitywas again measured using distilled water. After that, the set upwas ready for the next experiment with centrifuged stevia extract.All the experiments were conducted at a room temperature of32 ± 2 �C.

2.4.4. Nanofiltration experiments in stirred batch cellFirst, the membrane was compacted for 2 h at 800 kPa pressure

using distilled water. Water flux was measured at five differenttransmembrane pressure drop values. From the slope of the per-meate flux and pressure drop curve, the membrane permeabilitywas determined. Next, the cell was filled up by 300 ml of ultrafil-tered stevia extract (collected at 550 kPa and 100 l/h) and the oper-ating pressure was set using the nitrogen cylinder through aregulator. The stirring speed in the cell was set at an appropriaterpm by using a variac. Each experiment was conducted for 1 h atthe room temperature of 30 ± 2 �C. Clarified stevia extract (perme-ate) was collected in a measuring cylinder. Cumulative volume ofpermeate as a function of time was measured. From the slope ofthe cumulative volume–time plot, the permeate flux as a functionof time was obtained. At the end of the experiment, the permeatesamples were collected and analyzed for total solids, color, clarityand stevioside concentration. After the experiment, the set up wasdismantled and the membrane was rinsed with distilled watercarefully and was kept in 2% surfactant, sodium dodecyl sulfatesolution overnight. The cleaned membrane was rinsed carefullyso that the traces of surfactant were removed and again its perme-ability was measured using distilled water.

2.5. Experimental design

Cross flow ultrafiltration experiments under total recycle modewere performed using the operating pressure difference as 276,414, 552 and 690 kPa. The cross flow rates were 60, 80, 100 and120 l/h. These experiments under batch concentration mode wereundertaken at 276, 414 and 552 kPa pressure and 100 l/h crossflow rate.

The operating variables for nanofiltration experiments weretransmembrane pressure drop and the stirring speed. The operat-

ing pressure drops were 827, 965, 1103 and 1241 kPa. At eachpressure drop, three stirrer sppeds, namely, 500, 1000 and 1500were used.

2.6. Analysis

The original, centrifuged and ultrafiltered stevia extract wereanalyzed for their color, clarity, total solid content and steviosideconcentration. Color of the extract was measured in terms of opti-cal absorbance (A) at a wavelength of 420 nm using a spectropho-tometer (BIORAD Smart Spec 3000, USA). Clarity of the extract wasmeasured in terms of percentage of transmittance (%T) using aspectrophotometer. This is given by the equation, %T=100 � 10�A,where, A is optical absorbance at a wavelength of 660 nm.

Total solid of the sample was measured gravimetrically by heat-ing the extract in hot air oven at 104 ± 2 �C until the difference inthe weight of the extract becomes constant at successive intervals[29]. Total solid was represented in terms of gram per 100 ml ofstevia extract.

Amount of stevioside present in the stevia extract (before andafter clarification) was analyzed by HPLC (M/s, Perkin Elmer Co.,Shelton, Connecticut, USA). Initially a calibration curve was pre-pared using standard stevioside of 98% purity. Then, for each anal-ysis, a sample volume of 30 ll was injected into an Agilent ZorbaxSB C-18 reverse phase HPLC column (4.6 mm ID, 250 mm lengthand 5 lm particle size). Mobile phase was a mixture of acetonitrileand water at a ratio of 80:20 (volume basis) and the flow rate was1 ml/min. UV detector (Perkin Elmer series 200) was used fordetection of stevioside present in the sample at a wavelength of210 nm. Stevioside concentration was represented in terms of per-centage (%) of stevioside recovered in the clarified extract as:

%Stevioside ¼ ðstevioside retained in clarifiedextract=stevioside present in feedÞ � 100:

2.7. Statistical Analysis

All membrane filtration experiments were carried out in tripli-cate. All the property data presented in the tables are reported withstandard deviation. The permeate flux data were within ±3% vari-ation and these were presented as error bars in the figures. Thepermeability values had ±5% variation.

Table 1Various properties of the ultrafiltered liquor at different operating conditions under total recycle mode of operation.

Operatingpressure (kPa)

Flow rate in l/h (crossflow velocity in m/s)

Permeatecolor (A)

Permeateclarity (%T)

Permeate totalsolid (g/l)

Permeatestevioside (%)

Purity of steviosideStevper

TSper

� � Selectivity of

steviosideStevper

LMWper

� �

276 60 (0.10) 1.5 ± 0.13 53.1 ± 1.1 17 71.8 ± 3.1 0.67 2.080 (0.14) 1.4 ± 0.12 62.5 ± 1.3 15 58.0 ± 2.1 0.61 1.6

100 (0.17) 1.1 ± 0.13 69.2 ± 2.3 12 49.0 ± 2.4 0.64 1.8120 (0.20) 1.1 ± 0.15 69.5 ± 1.3 11 49.0±3.2 0.70 2.4

414 60 (0.10) 0.9 ± 0.14 80.9 ± 1.5 14 49.0 ± 3.5 0.55 1.280 (0.14) 0.9 ± 0.13 81.3 ± 2.1 14 45.0 ± 2.8 0.51 1.0

100 (0.17) 0.9 ± 0.12 81.8 ± 2.3 13 43.3 ± 2.6 0.52 1.1120 (0.20) 0.8 ± 0.14 81.5 ± 2.4 11 40.4 ± 2.9 0.58 1.4

552 60 (0.10) 0.8 ± 0.15 83.4 ± 1.5 13 45.0 ± 2.4 0.55 1.280 (0.14) 0.7 ± 0.16 84.9 ± 2.6 11 40.5 ± 3.2 0.58 1.4

100 (0.17) 0.7 ± 0.12 85.5 ± 2.4 10 39.7 ± 2.6 0.63 1.7120 (0.20) 0.6 ± 0.13 87.3 ± 1.8 10 37.3 ± 3.5 0.59 1.4

690 60 (0.10) 0.6 ± 0.14 85.3 ± 1.7 11 37.2 ± 2.6 0.53 1.180 (0.14) 0.6 ± 0.12 86.1 ± 1.6 11 29.7 ± 2.8 0.43 0.7

100 (0.17) 0.6 ± 0.15 86.9 ± 2.2 10 31.1 ± 2.5 0.49 1.0120 (0.20) 0.4 ± 0.12 87.5 ± 2.3 10 27.5 ± 2.2 0.43 0.8

Centrifuged extract (feed) _ 10.31 ± 1.13 1.3 ± 0.5 27 ± 1.3 15753 ± 5.2 (mg/l)Crude extract _ _ _ _ 17167 ± 8.4 (mg/l)

Chhaya et al. / Separation and Purification Technology 89 (2012) 125–134 129

3. Results and discussions

3.1. Cross flow ultrafiltration

As mentioned earlier, two modes were used for conductingcross flow ultrafiltration experiments. First total recycle mode,where, the permeate was recycled back and the feed concentrationwas maintained constant. Second, the batch concentration mode,where the feed concentration was not recycled and the volumeof the feed tank was continued to reduce and the feed concentra-tion increased.

The ultrafiltration feed essentially contains various compo-nents, of which only Stevioside is desirable. The mixture of compo-nents can be clubbed together as high molecular weightcomponents (HMW). Components with lower molecular weightthan Stevioside are grouped in low molecular weight solutes(LMW). It is considered that the HMW is completely rejected,LMW is freely permeable, and Stevioside is partially retained bygel layer and membrane. Following this categorization, one canessentially get an estimate of the amount of LMW in the permeate(which is equivalent to that present in feed). LMW in permeate canbe estimated by TSper�Stevper. With this definition, one can deter-mine the purity and selectivity of Stevioside in permeate by,

purity ¼ stev ioside concentration in permeateconcentration of total solids in permeate

and;

selectivity ¼ stev ioside concentration in permeateconcentration of LMW in permeate

:

These values are presented in Table 1 for different operatingconditions in total recycle mode and in Fig. 7 in case of batch mode.

3.1.1. Total recycle modeIn this mode of operation, variation of permeate flux as a func-

tion of time for various transmembrane pressure drop and thecross flow rates are shown in Fig. 4. Three general trends are ob-served from these figures. First, the permeate flux declines overtime of operation and finally, a steady state is reached. Second, atany point of time, the permeate flux increases with cross flow ratesat a fixed transmembrane pressure drop. Third, at any point oftime, the permeate flux increases with transmembrane pressuredrop at a fixed cross flow rate. The first observation is due to con-

centration polarization. As time of filtration progresses, more sol-utes are convected towards the membrane and a cake type oflayer starts growing over the membrane surface. Thickness of thislayer increases with the time of filtration. This layer offers a resis-tance against the solvent flux. Since, with time of filtration, thick-ness of this layer increases, the permeate flux declines. Forexample, at 276 kPa operating pressure and 120 l/h cross flow rate,the permeate flux decreases from about 18 to 14 l/m2 h (a decreaseof 22%) after 45 min (Fig. 4a). At the same transmembrane pressuredrop and 80 l/h cross flow rate, the decrease in flux is about 35%over 45 min (Fig. 4a). Similar, decline in flux over the filtrationduration for 414 kPa at 120 and 80 l/h of cross flow rates is 20%and 38% (Fig. 4b). These values of flux decline for 552 kPa are20% and 22% at these flow rates (Fig. 4c). At 690 kPa, these valuesare 17% and 16% (Fig. 4d). Therefore, the permeate flux declinesover the filtration time in between 16% and 38% for different trans-membrane pressure drop and the cross flow rates. Thus, at highercross flow rate, the flux decline is less. This is due to the fact that athigher cross flow rate, thickness of the cake type layer decreasesdue to increased forced convection. It is also observed from thesefigures that a steady state is attained in all the cases. Initially,the convective flux of solutes towards the membrane due to pres-sure gradient is more and more solutes are deposited over themembrane surface, forming a cake layer. This layer keeps on grow-ing as more solutes are convected towards the membrane. Aftersometime, the growth of this layer is arrested by the forced con-vection imposed by the cross flow rate in the flow channel and asteady state is attained. As observed from Fig. 4a, that at 276 kPapressure, steady state is attained after about 22 min. This time isreduced 10–20 min at higher operating pressure drops.

At a fixed transmembrane pressure drop, the permeate flux in-creases with the cross flow rates. At higher cross flow rate, theshearing action of the convective flow on the cake layer is moreand its growth is restricted. Therefore, the resistance against thesolvent flow offered by the cake layer is less, leading to an increasein permeate flux. For example, at 276 kPa pressure drop, the steadystate flux increases from 5 to 15 l/m2h as the cross flow rate in-creases from 60 to 120 l/h, affecting an increase of 200%. This in-crease is 125% at 414 kPa, 100% for 551 kPa and 83% at 690 kPa.

At a fixed cross flow rate, the permeate flux also increases withthe transmembrane pressure drop. Increase in pressure drop hastwo opposing effects. First, it increases the driving force leadingto flux enhancement. Second, more solutes are convected towards

0 10 20 30 40

6

8

10

12

14

16

18

20

22

24 276 kPa 60 l/h 80 l/h 100 l/h 120 l/h

Error bar: ±3%

Per

mea

te f

lux

( L

/m2 h)

Time (min)

0 10 20 30 4010

15

20

25

552 kPa 60 l/h 80 l/h 100 l/h 120 l/h

Error bar:±3%

Per

mea

te f

lux

(l/m

2 h)

Time (min)

0 10 20 30 40

15

20

25

30

35

Per

mea

te f

lux

(l/m

2 h)

Time (min)

690 kPa 60 l/h 80 l/h 100 l/h 120 l/h

Error bar:± 3%

0 10 20 30 40

10

15

20

25

30

35 414 kPa

60 l/h 80 l/h 100 l/h 120 l/h

Error bar:±3%

Per

mea

te f

lux

(l/m

2 h)

Time (min)

a b

c d

Fig. 4. Flux decline profiles during ultrafiltration in total recycle mode: (a) 276 kPa; (b) 414 kPa; (c) 552 kPa; (d) 690 kPa.

0 100 200 300 400 500 600 7000

3

6

9

12

15

18

21

24

27

30

Stea

dy s

tate

per

mea

te f

lux

(l/m

2 .h)

Operating pressure (kPa)

60 l/h 80 l/h 100 l/h 120 l/h

Error bar: ±3%

Fig. 5. Variation of steady state permeate flux with transmembrane pressure dropand cross flow rate.

0 1 2 3 4 5 6 7 8 9 10 110

1

2

3

4

5

6

7

8

9

10

11

12

Per

mea

te f

lux

(l/m

2 .h)

Time (h)

Flow rate = 100 l/h 276 kPa 414 kPa 552 kPa

Error bar: ±3%

1.00

1.05

1.10

1.15

1.20

1.25

1.30

1.35

1.40

Volum

e Concentration F

actor (VC

F)

Fig. 6. Flux decline profile and variation of volume concentration factor withtransmembrane pressure drop in batch concentration mode of cross flowultrafiltration.

130 Chhaya et al. / Separation and Purification Technology 89 (2012) 125–134

the membrane surface, thereby increasing the thickness of the cakelayer, resulting to increase in resistance against the solvent flowand consequently decline in permeate flux. However, by observingthe trends of flux decline profiles in Fig. 4a–d, it is clear that thefirst effect dominates the second one and the flux increases with

pressure. For example, at 60 l/h cross flow rate, the steady statepermeate flux increases from 5 to 12 l/m2 h (140% flux enhance-ment) as the pressure increases from 276 to 690 kPa. This enhance-ment over this pressure range is about 100% for 80 l/h, 81% for100 l/h and 57% for 120 l/h cross flow rate.

0 2 4 6 8 100.0

0.3

0.6

0.9

1.2

1.5

1.8

Col

or (

A)

Time (h)

100 l/h 276 kPa 414 kPa 552 kPa

0 2 4 6 8 10

60

80

100

120

Cla

rity

(%

T)

Time (h)

100 l/h 276 kPa 414 kPa 552 kPa

0 2 4 6 8 10

0.6

0.9

1.2

1.5

1.8

2.1

2.4

2.7

3.0

Tota

l sol

id (g

/100

mL

)

Time (h)

100 l/h 276 kPa 414 kPa 552 kPa

0 2 4 6 8 100

20

40

60

80

100

Stev

iosi

de R

ecov

ery

(%)

Time (h)

100 l/h 276 kPa 414 kPa 552 kPa

0 2 4 6 8 100.40

0.45

0.50

0.55

0.60

0.65

0.70

Pur

ity

Time (h)

276 kPa 414 kPa 552 kPa

a b

c d

e

Fig. 7. Profiles of permeate properties for various operating conditions in batch concentration mode of ultrafiltration: (a) color; (b) clarity; (c) total solids; (d) recovery ofstevioside; (e) purity of stevioside.

Chhaya et al. / Separation and Purification Technology 89 (2012) 125–134 131

The steady state flux values at different operating pressure dropand cross flow rates are presented in Fig. 5. The trends are as ex-pected and the reasons are already discussed earlier. The proper-ties of the permeate at the steady state with different operatingconditions are presented in Table 1. Some general trends are ob-served from this table. As the operating pressure drop increases,the stevioside recovery in the permeate decreases. The selectivity

and purity of stevioside in the permeate are almost independentof flow rate. Both selectivity and purity decrease with transmem-brane pressure drop. At higher pressure drop, the cake layer be-comes compact (associated with increasing porosity) and it actsas a dynamic membrane. Therefore, this layer retains some of thestevioside and recovery of stevioside in the permeate becomes less.For example, average (over various cross flow rates) recovery of

0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0 1.1 1.215

20

25

30

35

Per

mea

te f

lux

(l/m

2 .h)

Time (h)

Operating pressure = 827 kPa 500 rpm 1000 rpm 1500 rpm

Error bar: ± 3%

1.0

1.1

1.2

1.3

1.4

1.5

1.6

Vol

ume

Con

cent

rati

on F

acto

r (V

CF

)

0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0 1.115

20

25

30

35

40

45

Per

mea

te f

lux

(l/m

2 .h)

Time (h)

Operating pressure = 965 kPa 500 rpm 1000 rpm 1500 rpm

Error bar: ±3%

1.0

1.1

1.2

1.3

1.4

1.5

1.6

1.7

Volum

e Concentration F

actor (VC

F)

0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0 1.115

20

25

30

35

40

45

50

Per

mea

te f

lux

(l/m

2 .h)

Time (Hours)

Operating pressure = 1103 kPa 500 rpm 1000 rpm 1500 rpm

Error bars: ±3%

1.0

1.2

1.4

1.6

1.8 Volum

e Concentration F

actor (VC

F)

0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0 1.120

25

30

35

40

45

50

55

60

Per

mea

te f

lux

(l/m

2 .h)

Time (Hours)

Operating pressure = 1241 kPa 500 rpm 1000 rpm 1500 rpm

Error bar:±3%

1.0

1.2

1.4

1.6

1.8

2.0 Volum

e Concentration F

actor (VC

F)

a b

c d

Fig. 8. Flux decline profiles and variation of volume concentration factor with operating conditions during stirred batch nanofiltration: (a) 827 kPa; (b) 965 kPa; (c) 1103 kPa;(d) 1241 kPa.

132 Chhaya et al. / Separation and Purification Technology 89 (2012) 125–134

stevioside at 276 kPa is 56% and it is 44% at 414 kPa, 40% for552 kPa and 31% for 690 kPa. This dynamic cake type layer retainsother solids at higher pressure drop, thereby, increasing the clarityof the permeate remarkably at higher operating pressure. Clarity isabout 87% at 690 kPa and 120 l/h cross flow rate whereas, that inthe feed of ultrafiltration is only 1.26%. Thus, the total solids inthe permeate also decreases at higher operating pressure. It is alsonoted from this table that the stevioside recovery decreases mar-ginally with cross flow rates. Except the first two experiments, at276 kPa, 60 and 80 l/h, the variation of permeate recovery for dif-ferent cross rate at a fixed pressure value is insignificant. This isdue to the fact that the membrane for the first experiment is freshand after one experiment, there occurs some irreversible foulingthat reduces the stevioside recovery drastically. However, this foul-ing is present for subsequent experiments, but it is marginal andthe stevioside recovery shows a declining trend (although extre-mely small) with the cross flow rates at a fixed transmembranepressure drop.

3.1.2. Batch concentration modeAs mentioned earlier, in this mode of operation, the permeate

flux is not recycled back. In fact, for clarification of the stevia ex-tract, the permeate is the product and this is the most favorablemode of operation. Three experiments were conducted in this case,at 100 l/h cross flow rate, the operating pressure difference werevaried as 276, 414 and 552 kPa. The permeate flux profile alongwith the volume concentration ratio is presented in Fig. 6, as afunction of time. Two general trends are observed from this figure.

First, the permeate flux decline is more in this case compared tothe total recycle mode (in Fig. 4a–d) and second, there exists nosteady state. Flux decline is more at higher operating pressure. Inthis mode of operation, the permeate is not recycled to the feedtank; as a result, the volume of the feed tank reduces leading toan increase in feed concentration. As the feed concentration in-creases, the concentration polarization becomes more severe. Moresolutes are convected towards the membrane surface, resulting toa thicker cake layer. This increases the resistance against the sol-vent flux and the permeate flux declines. At higher operating pres-sure, solute deposition on the membrane surface is augmented byforced convection, leading to a further decline in permeate flux. Asthe above phenomena increases as a time of filtration increases, asteady state is never attained. Over a period of 10 h of operation,the flux decline is 6 to 2 l/m2 h at 276 kPa. It is 8.2 to 2.2 l/m2 hat 414 kPa and 11.5 to 3 l/m2 h at 552 kPa. Volume concentrationfactor (VCF) is defined as V0/V, where, V0 is the initial volume offeed and V is the volume at ant time. Variation of VCF with timeis presented in Fig. 6. It is observed from this figure that VCF ismore at higher pressure as more volume of permeate is filtrated.After 10 h of operation, VCF reaches a value of 1.35 at 552 kPa pres-sure and 100 l/h cross flow rate.

The properties of the permeate were also monitored over thefiltration period. Profiles of color, clarity, total solids, steviosiderecovery and purity in permeate are presented in Fig. 7a–e. It is ob-served from these figures that color, total solids and the steviosiderecovery in permeate decrease with time and clarity increases withtime. As discussed earlier, with progress in filtration time, the cake

Table 2Various properties of permeate of nanofiltration at the end of the experiment (feed is ultrafiltration permeate at 552 kPa and 100 l/h).

Operatingpressure (kPa)

Stirringspeed (rpm)

Permeatecolor (A)

Permeateclarity (%T)

Permeatetotal solid(g/100 ml)

Permeatestevioside(mg/l)

Product(retentate)total solid(g/100 ml)

Product (retentate)stevioside (mg/l)

Overall Recovery(%) [UF + NF]

Overall Purity(%)[UF + NF]

827 500 0.02 ± 0.003 99.5 ± 0.4 0.2 ± 0.05 127.9 ± 1.4 1.1 ± 0.05 6486 ± 5 41.3 59.01000 0.02 ± 0.002 99.9 ± 0.3 0.2 ± 0.04 143.1 ± 1.4 1.2 ± 0.04 7250 ± 3 46.2 60.41500 0.02 ± 0.003 99.1 ± 0.4 0.2 ± 0.02 196.8 ± 2.4 1.3 ± 0.02 7462 ± 6 47.5 57.4

965 500 0.02 ± 0.003 99.0 ± 0.5 0.2 ± 0.06 173.0 ± 3.4 1.2 ± 0.06 7283 ± 4 46.4 60.71000 0.02 ± 0.002 99.5 ± 0.4 0.2 ± 0.05 107.7 ± 2.4 1.3 ± 0.05 7606 ± 8 48.4 58.51500 0.02 ± 0.004 99.6 ± 0.3 0.2 ± 0.04 208.5 ± 1.4 1.3 ± 0.04 7976 ± 3 50.8 61.4

1103 500 0.01 ± 0.005 99.3 ± 0.2 0.2 ± 0.05 348.0 ± 3.4 1.3 ± 0.05 7565 ± 5 48.2 58.21000 0.02 ± 0.003 98.6 ± 0.3 0.2 ± 0.04 179.2 ± 2.4 1.4 ± 0.04 8233 ± 2 52.4 58.81500 0.02 ± 0.002 98.1 ± 0.2 0.2 ± 0.03 349.5 ± 3.4 1.4 ± 0.03 8529 ± 7 54.3 60.9

1241 500 0.02 ± 0.004 99.3 ± 0.4 0.1 ± 0.04 246.5 ± 2.4 1.4 ± 0.04 7952 ± 6 50.6 56.81000 0.02 ± 0.002 99.3 ± 0.4 0.1 ± 0.03 276.7 ± 1.4 1.5 ± 0.03 8726 ± 8 55.6 58.21500 0.02 ± 0.005 99.1 ± 0.3 0.2 ± 0.04 201.3 ± 2.4 1.6 ± 0.04 9594 ± 10 61.1 60.0

UF extract(feed forNF)

552 kPa, 100 l/h 0.62 ± 0.003 92.8 ± 0.3 0.9 ± 0.05 4945 ± 5.4

Centrifugeextract

Optimumoperatingcondition

10.6 ± 1.4 _ _ 14129 ± 6.2

Crude extract Optimumoperatingcondition

12.3 ± 1.6 0.006 ± 0.003 2.9 ± 0.4 15699 ± 4.4

Chhaya et al. / Separation and Purification Technology 89 (2012) 125–134 133

type of layer grows on membrane surface that acts as dynamicmembrane and retains the solutes. Thus, total solids and steviosiderecovery decreases. Although, color decreases with time, its varia-tion is marginal. An interesting observation is made from Fig. 7d.Stevioside recovery at the end of 10 h for all three operating pres-sures is in between 30% (at higher pressure, i.e. 552 kPa) and 38%(at lower pressure, i.e. 276 kPa). This is due to the enhanced reten-tion of the dynamic membrane at higher pressure by making itmore compact. From Fig. 6, it is also observed that the permeateflux after 10 h is 2 l/m2 h at 276 kPa that is marginally less thanthat at 552 kPa (3 l/m2 h). Since, stevioside recovery is our mainconcern, a lower transmembrane pressure drop must be selectedwith a reasonable permeate flux. On the other hand, the cross flowrate should be maximum to obtain a higher permeate flux (referFig. 4). Thus, among the operating conditions studied herein,276 kPa pressure and 120 l/h cross flow rate are suitable operatingconditions for ultrafiltration of stevia extract with 30 kDa mem-brane. Another interesting feature that can be observed fromFig. 7e is that the purity in case of lower pressure decreases onincreasing time of operation, which is not the case for higher pres-sure. The compactness of the dynamic cake layer over the mem-brane is not significant enough to screen other solids at lowerpressure. This results to increase in total solids in the permeate (re-fer to Fig. 7c), thereby, decreasing purity. It must be noted here,that purity does not exclusively depend on the amount of Stevio-side present in the permeate, rather, it the relative ratio of theamount of Stevioside to total solids in the permeate. So, even forthe same amount of Stevioside content, purity can be increases ifthe total solid content is decreased. So, ideally, the batch operationat lower pressure and high flowrate should be limited up to 5 h (asobserved from the present study) for maintaining high purity ofstevioside in the permeate.

3.2. Concentration by nanofiltration

Profiles of permeate flux and volume concentration factor withtransmembrane pressure drop are shown in Fig. 8a–d at variousstirring speeds. It is observed from these figures that the permeateflux declines with time and the flux is higher at higher operatingpressure, as expected. At 1241 kPa pressure and 1500 rpm, flux isthe highest and hence, volume concentration factor is the maxi-

mum to about 2 in the test cell. Various properties of the permeateare reported in Table 2. It is observed from this table that the clar-ity of permeate is more than 99% in most of the cases. Color andtotal solids in the permeate are quite low. Retention of color is inbetween 96% and 98% for different operating conditions studiedherein. Retention of stevioside is in the range of 93% to 98%. Theconcentration factor (ratio of feed concentration of stevioside toits initial concentration in feed) of stevioside is also presented inTable 2. It is observed from this table that at 1241 kPa pressureand 1500 rpm, the feed is concentrated about two times in 1 h ofoperation.

It can be concluded that the purity of overall process (UF + NF)is constant around 60%. However, the overall recovery of Steviosideincreases with stirring and transmembrane pressure drop. Maxi-mum recovery is obtained at 1241 kPa and 1500 rpm.

4. Conclusion

Clarification of centrifuged stevia extract by cross flow ultrafil-tration and subsequent concentration by nanofiltration was stud-ied in this work. In continuous cross flow under total recyclemode, permeate flux declined in between 16% and 38% for differentoperating conditions. However, steady state flux increased upto200% when the cross flow rate increased from 60 to 120 l/h at276 kPa. Flux enhancement upto 140% was attained when trans-membrane pressure drop increased from 276 to 690 kPa. At higheroperating pressure, recovery of stevioside in the permeate was less.Average 56% recovery of stevioside was obtained in the permeateat 276 kPa. However, in batch concentration mode of cross flowultrafiltration, about 38% stevioside recovery was attained after10 h. During nanofiltration, stevioside in the feed was concentratedtwice in 1 h at 1241 kPa pressure drop and 1500 rpm. Maximumrecovery is attained by ultrafiltration at 552 kPa, 100 l/h followedby nanofiltration at 1241 kPa and 1500 rpm.

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