Techno-economic evaluation of membrane cascades relative to simulated moving bed chromatography for...

Separation and Purification Technology 80 (2011) 600–609

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Separation and Purification Technology

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Techno-economic evaluation of membrane cascades relative to simulatedmoving bed chromatography for the purification of mono- and oligosaccharides

Johan Vanneste ⇑, Stijn De Ron, Steven Vandecruys, Sandra Adina Soare, Siavash Darvishmanesh,Bart Van der BruggenLaboratory of Applied Physical Chemistry and Environmental Technology, Department of Chemical Engineering, K.U. Leuven, Willem de Croylaan 46, B-3001 Leuven, Belgium

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

Article history:Received 24 March 2011Received in revised form 7 June 2011Accepted 9 June 2011Available online 24 June 2011

Keywords:Membrane cascadesSimulated moving bed chromatographySugars

1383-5866/$ - see front matter � 2011 Elsevier B.V. Adoi:10.1016/j.seppur.2011.06.016

⇑ Corresponding author. Tel.: +32 16 32 23 49; fax:E-mail address: johan.vanneste@cit.kuleuven.be (J

In this paper the McCabe–Thiele method, previously used for the design of gas separation membrane cas-cades, was adapted for membrane cascades for solute–solute separations. This method was applied hereon three different sugar separations: raffinose–sucrose, fructose–glucose and xylose–glucose. The state ofthe art for all these separations is simulated moving bed (SMB) chromatography. For all separations amembrane cascade could be designed to reach the same specifications as the reference SMB. This is espe-cially remarkable for the very challenging glucose–fructose separation where starting from 50% fructosepurity a 94% fructose purity could be reached. Due to the high number of required stages, for this sepa-ration the cascade cost was several times higher than the cost of the reference SMB. However, for the raf-finose–sucrose and glucose–xylose separation the cascade cost was similar or lower than the cost of theSMB. Moreover from the raffinose–sucrose separation and purification it could be concluded that thecompetitiveness of membrane cascades over SMB chromatography increases with the plant size. Alsoif the purity requirement becomes less stringent the competitiveness increases as could be seen fromthe glucose–xylose separation. As a result membrane cascades seem most promising for large scale con-tinuous processes for producing pure but depending on the selectivity not ultrapure products. A hybridmembrane cascade SMB process could be envisaged to also cover the ultrapure products range in a costeffective way.

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1. Introduction

Ultrafiltration (UF) and nanofiltration (NF) membranes havebeen tested extensively for the purification of mono- and oligosac-charide solutions. A large amount of retention and flux data isavailable in literature for different applications and conditions[1–10]. However, the most relevant measures of the separationperformance from an industrial point of view, namely the productpurity and yield, are generally not mentioned or simply insuffi-cient. For most products the purity is a constraint set by legislationfor safety reasons or to be marketable as a pure compound. Forfood applications purity requirements of 90% and more are verycommon. Yield requirements are less stringent but processes witha yield much lower than 70% can be considered as not economi-cally viable. Feng et al. [9] studied the separation of monosaccha-rides and lactose from galacto-oligosaccharides and obtained agalacto-oligosaccharide fraction with 54.5% purity and 70% yield.Li et al. [10] studied the separation of monosaccharides andsucrose from fructo-oligosaccharides and obtained a fructo-

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+32 16 32 29 91.. Vanneste).

oligosaccharide fraction with 90% purity and 50% yield. The lowpurity or yield that can be achieved in single-stage membrane pro-cesses is an important reason why simulated moving bed chroma-tography (SMB) is still the state of the art technology for industrialsugar purification [11–14].

Although the development of specialized membranes willfurther decrease the separation performance gap with SMB in thefuture, a new strategy to improve the selectivity of membraneprocesses for the separation of solutes has recently been adopted:the integration of membranes with moderate selectivity into mem-brane cascades. Lightfoot [15] was, to the best knowledge of theauthors, the first to apply cascade theory, that originated from iso-tope separation, to solute–solute separations with membranes.Lightfoot [16] and Gunderson et al. [17] calculated the yield andpurity for a three-stage diafiltration membrane cascade for theseparation of a-lactalbumin and b-lactoglobulin with a membranewith a high selectivity of 21. A major disadvantage of a diafiltrationmembrane cascade is, however, that it cannot be operated contin-uously despite the fact that normal membrane filtration is, in con-trast with normal chromatography, an inherently continuousprocess. Although the separation performance of a diafiltrationcascade could be higher than a normal membrane cascade and as

J. Vanneste et al. / Separation and Purification Technology 80 (2011) 600–609 601

a result less stages would be needed to attain a certain yield or pur-ity, the effective number of membrane stages can still be higher asa consequence of the need for concentration of the permeatestreams and solvent recovery between the stages by a reverseosmosis step. The four-stage ultrafiltration cascade proposed byMayani et al. [18] is in fact a two-stage cascade with two stagesfor product concentration and generation of diafiltration buffer.Because of the simpler implementation in practice and the contin-uous operation, this study will focus on membrane cascades with-out diafiltration. Caus et al. [19] already explored the separationperformance of countercurrent membrane cascades without diafil-tration up to four stages for the separation of maltose and xylose.An elaborate design methodology to meet given separationrequirements was, however, still lacking. In this paper the McCa-be–Thiele method, used for the design of gas separation membranecascades [20], will be adapted for solute separation with mem-brane cascades. The membrane cascades will be optimized forminimal total membrane area and pumping duty. This will be doneby minimization of mixing losses as it is known from isotope sep-aration cascades that minimization of mixing losses automaticallyleads to minimal total interstage flows and in this case thusminimal total membrane area and pumping duty [21]. An algo-rithm was developed that returns for given single-stage membraneretentions and flux, a given feed purity and concentration and lastbut not least a given productivity, yield and purity requirement,the optimized membrane cascade configuration. As a result withthis algorithm a membrane cascade can be designed to do the samework of separation, expressed as the purity, the yield and the pro-ductivity, as a reference SMB which is absolutely essential to makea fair cost comparison of the different separation techniques.

Three applications were selected on the basis of the ease ofseparation for single-stage membrane processes. In a first casestudy the cost of a SMB process for the separation of raffinose(MW 504 g/mol) and sucrose (MW 342 g/mol) from sugar beetmolasses is compared with the cost of the equivalent membranecascade. Raffinose is concentrated in the molasses during theproduction process of sugar and acts as a crystallization inhibitorfor the remaining sucrose in the molasses. Each year 12 millionton of molasses are produced worldwide and the sucrose lossto the molasses can amount to 15% [22]. But more importantlyraffinose also has a beneficial effect on the growth of bifidobac-teria and constitutes therefore a high-value food additive [23].To study the effect of the scale on the costs for this separationthe costs will be calculated for both a pilot and an industrialscale plant. In a second case study a much more challenging sep-aration will be studied: the separation of a mixture of glucose(MW 180 g/mol) and fructose (MW 180 g/mol). Since the indexof sweetness of fructose is about twice as much as that of glu-cose and the glycemic index is four times lower, the conversionand separation of fructose from glucose is of great commercialimportance. The glucose–fructose mixture results from the enzy-matic conversion of 100% glucose in corn starch [24]. In order toproduce high fructose corn syrup (HFCS) containing 90% fructose,the fructose should be removed continuously as the conversionrate decreases sharply from about 50% fructose. Each year15 million tons of HFCS is produced with SMB worldwide. Thelast separation that will be studied is the separation of xylose(MW 150 g/mol) from glucose (MW 180 g/mol). A mixture ofglucose–xylose is obtained from hydrolysis of certain lignocellu-losic biomass. Xylose can be converted biochemically to xylitol.Xylitol is a natural insulin stabilizer that is beneficial for diabe-tes patients. Xylitol is also a non-fermentable sugar alcohol andas a result does not participate in the formation of dental caries.Glucose can inhibit the conversion to xylitol and should thus beremoved [4]. For this separation also the effect of the target pur-ity on the plant cost was studied.

2. Theory and methodology

2.1. Cascade theory

In 1975 Hwang and Kammermeyer [25] suggested to use theMcCabe–Thiele method, originally developed for distillation, to de-sign membrane cascades for the separation of binary gas mixtures.A McCabe–Thiele diagram is constructed from three equations: anequilibrium line equation, an operating line equation for the strip-ping section and an operating line equation for the enriching sec-tion. Starting from the feed composition the operating line of theenriching section is stepped off against the equilibrium line untilthe target purity is reached and the operating line of the strippingsection is stepped off against the equilibrium line until the targetyield is reached. The total number of steps on the McCabe–Thielediagram corresponds to the required number of stages to attain acertain purity and yield.

The equilibrium line equation gives the relation between themole fraction of the desired component 1 in the enriched fraction,y1, and in the depleted fraction, x1, of a stage. The equation is basi-cally a reorganization of the definition of the overall separationfactor a:

a ¼y1

1�y1x1

1�x1

ð1Þ

In distillation the equilibrium curve is determined by the ther-modynamic vapor–liquid equilibrium, while for membrane pro-cesses the separation characteristics of the module are ratedetermined. Although the term equilibrium line equation is thusinappropriate for membrane processes, it is used frequently andwill also be used here. The mole fractions, yi and xi, can also be cal-culated for membrane processes for solute–solute separations. Ifthe desired component 1 is in the permeate:

y1 ¼Q pcp1

Q pcp1 þ Q pcp2ð2Þ

Qp is the flow rate of the permeate in m3/s and cpi the concen-tration of compound i in the permeate in mole/m3. In this paperthe McCabe–Thiele diagram is always constructed as a functionof the mole fraction of the more permeable compound. Integratingthe definition of the retention of the membrane module with cfi theconcentration of compound i in the feed in mole/m3:

R ¼ 1� cp1

cf1ð3Þ

and the feed purity:

Pf1 ¼cf1

cf1 þ cf2ð4Þ

into Eq. (2) yields:

y1 ¼1

1þ ð1�R2Þð1�R1Þ

ðP�1f1 � 1Þ

ð5Þ

An analogous derivation can be made for the concentrate molefraction x1 and yields:

x1 ¼1

1þ CF2ð ÞCF1ð Þ P�1

f1 � 1� � ð6Þ

CFi is the concentration factor and is function of the membraneretention, Rmi, and the recovery, REC = Qp/Qf, of the membranemodule [26]:

CFi ¼1

ð1� RECÞRmið7Þ

602 J. Vanneste et al. / Separation and Purification Technology 80 (2011) 600–609

Filling in Eqs. (5) and (6) into Eq. (1) gives the overall separationfactor for membrane processes for solute–solute separations:

a ¼ 1� R1

1� R2

� �CF2

CF1

� �ð8Þ

The first factor of the separation factor corresponds to the selec-tivity of the membrane. From Eq. (8) it can be derived that theequilibrium line equation depends on the solvent recovery of themembrane module. Unlike for distillation, the equilibrium line willbe different for every stage if the recovery of the modules are notthe same.

The operating line equations give the relation between thecomposition of streams of subsequent stages in a cascade. Theycan be derived from total material and component balanceswhich are determined by the cascade configuration and operat-ing conditions. The most general countercurrent cascade is a cas-cade that sends the permeate of one stage to the feed of anystage higher in the cascade and sends the concentrate back tothe feed of any stage lower in the cascade. The simplest counter-current cascade is a so-called ‘one-up, one-down’ cascade. In thiscascade the permeate of one stage is sent to the next stage andthe concentrate is sent to the previous stage in the cascade.From an energetic point of view this type of cascade will prob-ably also be the most efficient one. Because the difference incomposition of the streams of stages will increase as the dis-tance between the stages increases, mixing losses will be mini-mal for a ‘one-up, one-down’ configuration [27]. As the cascadeoperating conditions will be optimized for minimal mixinglosses, it is logical that this study is confined to ‘one-up, one-down’ cascade configurations. For this type of cascade the oper-ating line equation of the enriching section with N stages and aproduct purity in the product stream yp is [25]:

ynþ1 ¼

PNi¼1

QNj¼i

cj

!

PNi¼1

QNj¼i

cj

!þ 1

xn þ1

PNi¼1

QNj¼i

cj

!þ 1

yp ð9Þ

The operating line equation for the stripping section with Mstages and a product purity in the waste stream xw is as follows[25]:

ym ¼1þ

PMi¼m

QMj¼1

1cj

!

PMi¼m

QMj¼1

1cj

! � xmþ1 þ1

PMi¼m

QMj¼1

1cj

! � xw ð10Þ

Both Eqs. (9) and (10) are a function of the stage cut hj of everystage because of [25]:

cj ¼1� hj

hjð11Þ

In order to be able to use Eqs. (9) and (10) for solute–solute sep-arations with membranes a relation should be derived between thestage cut h and REC, the recovery of a stage. The stage cut h is de-fined as [25]:

h ¼ np1 þ np2

nf1 þ nf2ð12Þ

With npi and nfi the number of moles of component i in perme-ate and feed, respectively. The recovery of a membrane module isdefined as:

REC ¼Q p

Q fð13Þ

Writing Eq. (12) as a function of concentrations yields followingequation:

h ¼c1fð1� R1ÞQ p þ c2fð1� R2ÞQ p

c1f Q f þ c2f Q fð14Þ

Incorporating the definition of the recovery (13) and the feedpurity (4) into Eq. (14) gives the final relation for the feed stage be-tween the stage cut h and the recovery REC.

h ¼ REC:Pf1:ðð1� R1Þ þ ðP�1f1 � 1Þð1� R2ÞÞ ð15Þ

There are many possible cascade configurations that can attainthe desired purity and yield. An optimization procedure can selectthe most appropriate configuration and operating conditions. Aninteresting objective function to minimize is the mixing losses asthis automatically leads to minimal total interstage flows and thusminimal total membrane area and pumping duty [21]. The increasein entropy DS due to mixing of two streams F1 and F2 can be cal-culated as follows [21]:

DS ¼ �RF½xF ln xF þ ð1� xFÞ lnð1� xFÞ��RF1½xF1 ln xF1 þ ð1� xF1Þ lnð1� xF1Þ��RF2½xF2 ln xF2 þ ð1� xF2Þ lnð1� xF2Þ�

ð16Þ

where x is the mole fraction of the desired component in everystream and R the gas constant. F is the total mole in a stream. In thisstudy the recovery was assumed constant over the stripping andenriching section but was allowed to differ between both sections.In order to limit fouling a maximum allowable module recovery of85% was chosen. Once the optimal configuration is selected, thealgorithm calculates all the streams on the basis of the feed and out-put concentrations and the optimized module recoveries. Eventu-ally the required membrane area of a stage can be calculated asthe ratio of the permeate flow in that stage to the single-stagemembrane flux.

2.2. Cost comparison method and assumptions

Major SMB costs that have been identified are: the resin cost,the column cost, the pump cost, the cost for pumping and some-times also heating and the cost for evaporation to obtain a dryproduct. The cost for the eluent was not taken into account as thisis mostly water for sugar purification. To estimate the total resincost (US$) the following correlation can be used [28]:

Cres ¼ 3623000ðdresÞ�1:675mres ð17Þ

where dres is the diameter of the resin in lm and mres is the totalmass of resin in kg. If the total mass of resin is not given it can beestimated from [28]:

mres ¼ VcolNcolð1� eÞqres ð18Þ

With Vcol and Ncol the volume and number of columns, respec-tively. The packing void fraction e is usually around 40%. For thedensity of the resin qres a value of 800 kg/m3 was assumed. Thenumber and dimensions of the columns, more specifically thediameter dcol and height hcol, together with the operating pressurepcol (in kPa) determine the cost for the pressure vessel [29]:

Ctanks ¼ 583:6d0:675col hcolNcolFmatð

14:5pcol

50Þ0:44 ð19Þ

For food applications the standard material is stainless steel. Forthe AISI 304 type Fmat is equal to 1.7. If the operating pressure wasnot mentioned a value of 500 kPa was assumed. Apart from thecost for the pressure vessel, column costs also comprise a costfor the structured packing that keeps the resin in place [29]:

J. Vanneste et al. / Separation and Purification Technology 80 (2011) 600–609 603

Cpacking ¼pd2

col

4hcolCpacking specific ð20Þ

The specific packing cost depends on the material and the typeof packing. Here an AISI 316 stainless steel structured packing waschosen with a specific cost of 2374.7 €/m3. The total column costcan then be calculated as the sum of the costs for the tanks andthe packing:

Ccol ¼ Ctanks þ Cpacking ð21Þ

The costs for the pumps (US$) for the feed and eluent can beestimated from the pressure p (kPa) and the flow Q (m3/h) with fol-lowing correlation [30]:

Cpump ¼ 81:28f 1f2LðpQÞ0:39 ð22Þ

L is equal to 1.4 and is a factor to incorporate labor costs. Theconstruction material is AISI 304 for which f1 = 1.5. Suction heightis negligible so f2 = 1.0.

The resin cost, the column cost and the pump cost are consid-ered to be the most important capital costs for SMB. Major costsfor membrane cascades are similar as for SMB except that the resincost is replaced by the membrane cost. The membrane cost can becalculated from the total membrane area Amem as follows [30]:

Cmem ¼ Cmem specificAmem ð23Þ

For the specific membrane cost a cost of 100 €/m2 was assumed.The membrane area can be calculated as the ratio of the sum of thepermeate flows of all stages, which is calculated by the algorithm,to the single-stage membrane flux, which is obtained from litera-ture. The equivalent of the column cost is the membrane modulecost. The membrane module cost (US$) can also be estimated fromthe total membrane area [30]:

Cmod ¼ 3047:21ðAmemÞ0:53 ð24Þ

The pump costs of a membrane cascade can be calculated withEq. (22).

All the remaining costs are operating costs. The pumping, heat-ing and evaporation costs can be calculated in the same way forSMB as for membranes. The pumping costs (€/h) can be calculatedfrom the operating pressure p (Pa) and flow Q (m3/h) [30]:

Cpumping ¼1

3:6� 106

pQCe

gpumpð25Þ

A pump efficiency g of 70% was assumed. For electricity a costCe of 0.12 €/kW h was used [31]. The cost for heating (€/h) is deter-mined by the feed flow Q (m3/h) and the temperature T (�C) [32]:

Cheating ¼1

109

qsQCpsðT � 20ÞCg

gboilerð26Þ

The density qs and heat capacity Cps of the solution were calcu-lated from the feed composition. For the density of the sugars a va-lue of 1540 kg/m3 was used except for raffinose the density is equalto 1810 kg/m3. The heat capacity of the sugars was assumed to beequal to 1256 J/kg K. The density and heat capacity of water is1000 kg/m3 and 4186 J/kg K, respectively. A gas boiler efficiencyg of 90% was assumed. A gas price Cg of 9 €/GJ was used [33].The cost for evaporation (€/h) with a multi-effect evaporator canbe calculated from the total water flow Qw (m3/h) that has to beevaporated [34]:

Cevaporation ¼1

109

QwqwEsteamCg

GRgboilerð27Þ

The evaporation energy of steam is 2.2 MJ/kg. A gain ratio GR of5.7 was assumed.

In order to be able to compare the costs of SMB with the equiv-alent membrane cascade all the costs were annualized. Thus whena cost is mentioned further in the discussion, the annualized cost isconsidered. To annualize capital costs they should be multiplied bythe capital recovery factor A/P [35]:

A=P ¼ ið1þ iÞDL

ð1þ iÞDL � 1ð28Þ

A plant design life DL of 20 years and a discount rate i of 8%were assumed. In the case a capital cost P occurs in the future inyear t, first the net present value NPV was calculated and then mul-tiplied by the capital recovery factor [35]:

NPV ¼ Pð1þ iÞt

ð29Þ

This will be the case for the resin and the membrane cost. A de-sign life of both the resin and the membrane of 5 years was as-sumed. This means the resin and membrane cost occur fourtimes over the plant design life. This seems low but the treatedstreams are all sugar solutions which have a low fouling potential.Sugars dissolve well into water and as a result don’t have a ten-dency to attach to the membrane. The lifetime of the other compo-nents was assumed to be equal to the plant design life. Toannualize operating costs (€/h) they should be multiplied by theaverage operating hours per year. It was assumed that the plantoperates for 5600 hours per year. This corresponds to a standstillof a bit more than 4 months. Two to three months of which isdue to shortage of qualitative feedstock material and the remainingtime for maintenance. In European sugar refineries the campaignstarts typically end of September and ends end of June. Cost corre-lations where updated by multiplying with the ratio I of the Chem-ical Engineering Plant Cost Index (CEPCI) which can be retrieved inperiodicals such as Chemical Engineering. The column cost and theresin cost were multiplied by 1.36 and 1.38, respectively. Thepump cost and module cost were multiplied by 3.34 and 1.41,respectively. In case the costs were given in US$, they weremultiplied by 0.73 to convert into €.

The labor cost was only taken into account in the case of thepump costs. However, both SMB and industrial membrane pro-cesses are highly automated processes. Even if labor costs are notnegligible they can be considered similar for both techniques.The same assumption is plausible for the maintenance costs asthe maintenance operation with by far the highest cost, namelyreplacement of the resin for SMB and replacement of the mem-branes for the membrane cascade, has approximately the same fre-quency of once every 5 years for both techniques. Althoughcolumns can have large dimensions, the higher building cost dueto the higher footprint was not taken into account.

3. Results and discussion

3.1. Raffinose–sucrose separation

Raffinose differs from sucrose by a complete galactose mole-cule. The difference in retention of sucrose and raffinose and theselectivity of the membranes found in literature is, however, ratherlow. For the UFCA membrane the selectivity is 1.9 and the differ-ence in retention is 0.26 (Table 1). The UFCA membrane is a tightultrafiltration membrane with a molecular weight cut-off of1000 Da. For the NFCA membrane the selectivity is slightly higher,namely 2.2, and the difference in retention slightly lower, namely0.24. The NFCA membrane is a nanofiltration membrane with aNaCl retention of 50%. The feed stream of the pilot scale SMB con-tains 23 wt.% raffinose and attains a raffinose purity and yield of81.2% and 70.5%, respectively (Table 1).

Table 1Design specifications and characteristics for the separation of raffinose from sucrose with SMB and an equivalent membrane cascade with two differentmembranes NFCA and UFCA [5,36] (pilot scale).

SMB Equivalent membrane cascade NFCA UFCA

Raffinose feed [wt.%] 23 Raffinose feed [wt.%] 23 23Solids feed [wt.%] 60 Solids feed [kg/m3] 53.3 53.3Raffinose purity [–] 0.81 Raffinose purity [–] 0.83 0.90Raffinose yield [–] 0.71 Raffinose yield [–] 0.76 0.67Productivity [kg raffinose/year] 1956 Productivity [kg raffinose/year] 1956 1956

Retention [sucrose/raffinose] 0.56/0.8 0.45/0.71Desorbant consumption [liter water/kg raffinose] 6.2 Membrane flux [10�6m3/m2s] 5.0 5.5Diameter of resin [lm] 150 Membrane area [m2] 335 948Volume of resin [m3] 0.1106 Stages enriching [–] 2 3Number of columns [–] 8 Recovery enriching stage [–] 0.08 0.18Diameter of columns [m] 0.1 Stages stripping [–] 4 4Length of columns [m] 1.5 Recovery stripping stage [–] 0.83 0.81Feed rate [m3/s] 0.00044 Feed rate [m3/s] 0.00001 0.00001Eluent rate [m3/s] 0.0015Pressure [bar] 5 Pressure [bar] 6.9 6.9Temperature [�C] 70 Temperature [�C] 20 20Solids output [wt.%] 11.5 Solids output [kg/m3] 67 66

604 J. Vanneste et al. / Separation and Purification Technology 80 (2011) 600–609

The low starting purity and the low selectivity and difference inretention of the NFCA membrane are the main reasons why amembrane cascade with six stages is necessary to achieve the samepurity and yield as the reference SMB (Table 1). For the UFCA mem-brane, which has a lower selectivity, seven stages are needed to ob-tain the desired raffinose purity (Table 1). Due to the constraint tohave the same productivity, the additional stage leads to a dra-matic increase in membrane area from 335 to 948 m2 for the NFCAand UFCA membrane, respectively (Table 1). Fig. 1 gives theMcCabe–Thiele diagram for the NFCA membrane cascade. The dia-gram is constructed as a function of the mole fraction of the mostpermeable compound, in this case sucrose. The two top stages con-stitute the stripping section to obtain the desired yield for raffinoseor equivalently the enriching section to obtain the desired purityfor sucrose. The four bottom stages are needed to achieve the de-sired raffinose purity.

Although six membrane stages seems a lot, a membrane cas-cade with the NFCA membrane is affordable in comparison withthe reference SMB (Fig. 2). The cost for the SMB is 22,303 € whilethe cost for the membrane cascade with NFCA is equal to22,332 €. By far the major cost for SMB is the resin cost (83%).Changing this cost, by for instance using a resin with a higherdiameter, will thus have the largest impact on the total cost ofSMB. Changing the diameter of the resin will, however, change

Fig. 1. McCabe–Thiele diagram for NFCA membrane cascade for sucrose (pilot scale,red = stripping section, blue = enriching section, upper dashed line = equilibriumline) (For interpretation of the references to color in this figure legend, the reader isreferred to the web version of this article.).

the separation characteristics and thus the design. To study the ef-fect of a resin diameter change on the total costs, an SMB modelshould be developed which was not the aim of this study. The sec-ond most important SMB cost is the cost for the columns (15%). Thepump cost together with the operating costs (pumping, heating,evaporation) constitute the remaining 2% of the costs. The costsof the NFCA membrane process are more evenly spread. The largestcost is the cost for the membranes (38%). The second largest cost isthe cost of the modules (31%) followed by the pump cost (17%) andthe pumping cost (13%). The cost for evaporation constitutes theremaining 1% of the total costs as no heating was applied.

With 54,926 € the cost of the UFCA cascade is almost 2.5 timeshigher than the cost of the NFCA cascade despite the fact that onlyone extra stage is needed. Although the purity of the UFCA cascadewas higher, the yield was lower (Table 1). The much higher costsfor the UFCA cascade suggests that the costs of membrane cascadesincrease more than linearly with the number of stages. As the feedflow rate per stage is maximal at the feed stage in a cascade, all theflows in a cascade will increase on adding an extra stage to assurethe same productivity. Therefore, not only the cost of an extrastage is incurred but also the cost of the increased equipment sizein the other stages. That the flows increase sharply with the num-ber of stages is confirmed by the required membrane area, which isalmost three times higher for the UFCA membrane compared with

Fig. 2. Comparison of costs for the separation of raffinose from sucrose with SMBand an equivalent membrane cascade with two different membranes NFCA andUFCA (pilot scale).

Table 2Design specifications and characteristics for the separation of raffinose from sucrose with SMB and an equivalent membrane cascade withtwo different membranes NFCA and UFCA [5,23] (industrial scale).

SMB Equivalent membrane cascade NFCA

Raffinose feed [wt.%] 19 Raffinose feed [wt.%] 19Solids feed [wt.%] 60 Solids feed [kg/m3] 53.3Raffinose purity [–] 0.70 Raffinose purity [–] 0.84Raffinose yield [–] 0.75 Raffinose yield [–] 0.73Productivity [ton raffinose/year] 464 Productivity [kg raffinose/year] 464

Retention [sucrose/raffinose] 0.56/0.8Desorbant consumption [liter water/kg raffinose] 5.9 Membrane flux [10�6 m3/m2 s] 5.0Diameter of resin [lm] 150 Membrane area [m2] 100320Volume of resin [m3] 40 Stages enriching [–] 2Number of columns [–] 12 Recovery enriching stage [–] 0.08Diameter of columns [m] 1.681 Stages stripping [–] 4Length of columns [m] 1.5 Recovery stripping stage [–] 0.84Feed rate [m3/s] 0.006 Feed rate [m3/s] 0.003Eluent rate [m3/s] 0.022Pressure [bar] 5 Pressure [bar] 6.9Temperature [C�] 70 Temperature [C�] 20Solids output [wt.%] 11.5 Solids output [kg/m3] 67

J. Vanneste et al. / Separation and Purification Technology 80 (2011) 600–609 605

the NFCA membrane while the flux is approximately the same(Table 1).

This is really important because a small increase in membraneselectivity can thus lead to much lower costs of the membrane pro-cess and in this case a membrane cascade could become morecompetitive than SMB. An important prerequisite for a large costreduction by using a high selectivity membrane is of course thatthe flux of the membrane does not alter too much. Yet the develop-ment of membranes that combine a high permeability with a highselectivity does not seem to be an easy task.

The installation described in Table 1 with a productivity of al-most 2 ton per year is a pilot scale installation. In literature anindustrial SMB installation was described with a productivity of464 ton per year [23]. It is now possible to study the effect of theplant size on the cost as the industrial scale installation has a sim-ilar feed and end purity and yield as the pilot plant (Tables 1and 2). The amount of resin or membrane surface required scalesroughly linearly with the feed rate. Because the flow velocity isfixed to an optimal value a higher flow rate (=flow velocity � chan-nel cross section) will lead to more membrane modules in parallelor larger column diameters. Due to the lack of appropriate correla-tions incorporating the economies of scale into the resin and themembrane cost, these costs were assumed to be constant and asa result also evolve linearly with the feed rate. All the other costs

Fig. 3. Comparison of costs for the separation of raffinose from sucrose with SMBand the NFCA membrane cascade (industrial scale).

increase less than linearly with the feed rate. As the resin cost ac-counted for 83% of the total cost for the SMB and the membranecost accounted for only 38% of the total cost of the NFCA mem-brane cascade for the pilot plant, it can be expected that the costof the SMB will increase much more for the industrial installationthan the cost of the NFCA cascade.

From Fig. 3 it is clear that the cost of the NFCA membrane cas-cade (3.6 M€) is now almost half of the cost of the SMB installation(6.8 M€). The resin cost increased from 83% to 98% of the total SMBcost while the membrane cost increased from 38% to 69% of the to-tal cascade cost.

3.2. Glucose–fructose separation

Fructose and glucose both have the same molecular weight.They only differ by structure. This separation can be expected tobe much more difficult with membrane processes. This is also con-firmed by the difference in retention of the two compounds: 0.16for the NFCA membrane and only 0.10 for the UFCA membrane(Table 3). In this case only for a lab scale SMB enough technicaldata were found in literature to calculate the costs. The purity ofthe feed (50% fructose) is now much higher than with the raffi-nose–sucrose separation but also the yield (91%) and purity(95%) requirements are much more stringent. Together with thelow selectivity this leads to an NFCA cascade with 21 stages (Fig. 4).

From Fig. 5 it is clear that the almost two times smaller differ-ence in retention of the UFCA membrane leads to a much highercascade cost. The number of required stages increases almosttwo times from 21 to 38, but the cost of the UFCA cascade is almost10 times higher. This confirms again the apparently more than lin-ear relationship between the number of stages and the total cas-cade cost.

The high number of stages in combination with the low produc-tivity makes that the membrane cost is not the largest cost like inthe previous application. In fact for the NFCA membrane, the mem-brane cost is only 5% of the total cost. By far the largest cost now isthe pump cost (63%) followed by the module cost (18%) and thepumping cost (14%). For the same reason of low productivity theresin cost of the SMB is now responsible for only 11% of the costswhile the column cost is now the largest cost (76%, Fig. 6). In thirdplace comes the cost for the pumps (9%).Unfortunately the costs ofthe NFCA membrane cascade are more than 17 times larger thanthe costs of the SMB (Fig. 6). Although the separation can beachieved with a cascade consisting of commercial membranes,

Table 3Design specifications and characteristics for the separation of fructose from glucose with SMB and an equivalent membrane cascade with two differentmembranes NFCA and UFCA [5,37].

SMB Equivalent membrane cascade NFCA UFCA

Fructose feed [wt.%] 50 Fructose feed [wt.%] 50 50Solids feed [kg/m3] 80 Solids feed [kg/m3] 50.7 50.7Fructose purity [–] 0.95 Fructose purity [–] 0.96 0.95Fructose yield [–] 0.91 Fructose yield [–] 0.87 0.95Productivity [kg fructose/year] 41 Productivity [kg fructose/year] 41 41

Retention [fructose/glucose] 0.26/0.43 0.22/0.32Desorbant consumption [liter water/kg fructose] 109 Membrane flux [10�6 m3/m2 s] 10 12Diameter of resin [lm] 320 Membrane area [m2] 18 412

Stages enriching [–] 11 19Number of columns [–] 12 Recovery enriching stage [–] 0.36 0.38Diameter of columns [m] 0.026 Stages stripping [–] 10 19Length of columns [m] 0.3 Recovery stripping stage [–] 0.65 0.62Feed rate [m3/h] 0.0002 Feed rate [m3/h] 0.00033 0.0003Eluent rate [m3/h] 0.0008Pressure [bar] 5 Pressure [bar] 13.8 13.8Temperature [C�] 60 Temperature [C�] 20 20Solids output [kg/m3] 14 Solids output [kg/m3] 51 51

Fig. 4. McCabe–Thiele diagram for NFCA membrane cascade for fructose(red = stripping section, blue = enriching section, upper dashed line = equilibriumline) (For interpretation of the references to color in this figure legend, the reader isreferred to the web version of this article.).

Fig. 5. Comparison of costs for the separation of fructose from glucose with anequivalent membrane cascade with two different membranes NFCA and UFCA.

Fig. 6. Comparison of costs for the separation of fructose from glucose with SMBand an equivalent membrane cascade with the NFCA membrane.

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for this application a membrane with a much higher selectivity isnecessary in order to compete with SMB.

3.3. Glucose–xylose separation

The molar mass of a xylose molecule is 150 g/mol and thus dif-fers only slightly from that of a glucose molecule (180 g/mol). Inliterature the NF270 membrane was found to have a selectivityof 1.9 and a difference in retention of 0.34 (Table 4). The NF270membrane is a nanofiltration membrane with a MgSO4 retentionof 97%. A pilot scale SMB installation was found in literature withtwo operation modes: moderate purity and high yield on one hand(low purity mode, Table 4) and very high purity and moderateyield on the other hand (high purity mode, Table 5). In this waythe effect of the purity requirement on the cost can be evaluated.In the first operation mode, the feed contains already 67% xyloseand should be purified up to 94% xylose which is already ratherhigh. With the NF270 membrane this can be achieved with sixmembrane stages (Fig. 7).

In this first case of low purity the NF270 cascade is clearly morecost efficient than the SMB (�26%, Fig. 8). Due to the small scale

Table 4Design specifications and characteristics for the separation of glucose and xylose with SMB and an equivalent membrane cascade with NF270 membrane[4,38] (low purity).

SMB Equivalent membrane cascade NF270

Xylose feed [wt.%] 67 Xylose feed [wt.%] 67Solids feed [wt.%] 58 Solids feed [kg/m3] 335Xylose purity [–] 0.94 Xylose purity [–] 0.94Xylose yield [–] 0.86 Xylose yield [–] 0.76Productivity [kg xylose/year] 423 Productivity [kg xylose/year] 423

Retention [xylose/glucose] 0.28/0.62Desorbant consumption [liter water/kg fructose] 1.41 Membrane flux [10�6 m3/m2 s] 10Diameter of resin [lm] 350 Membrane area [m2] 0.25

Stages enriching [–] 3Number of columns [–] 20 Recovery enriching stage [–] 0.32Diameter of columns [m] 0.03 Stages stripping [–] 3Length of columns [m] 0.69 Recovery stripping stage [–] 0.70Feed rate [m3/h] 0.0003 Feed rate [m3/h] 0.00029Eluent rate [m3/h] 0.0010Pressure [bar] 5 Pressure [bar] 40Temperature [C�] 50 Temperature [C�] 50Solids output [kg/m3] 10.9 Solids output [kg/m3] 97

Table 5Design specifications and characteristics for the separation of glucose and xylose with SMB and an equivalent membrane cascade with NF270 membrane[4,38] (high purity).

SMB Equivalent membrane cascade NF270

Xylose feed [wt.%] 71 Xylose feed [wt.%] 71Solids feed [wt.%] 44 Solids feed [kg/m3] 335Xylose purity [–] 0.999 Xylose purity [–] 0.994Xylose yield [–] 0.79 Xylose yield [–] 0.727Productivity [kg xylose/year] 423 Productivity [kg xylose/year] 423

Retention [xylose/glucose] 0.28/0.62Desorbant consumption [liter water/kg fructose] 3.13 Membrane flux [10�6 m3/m2 s] 10Diameter of resin [lm] 350 Membrane area [m2] 0.28

Stages enriching [–] 6Number of columns [–] 20 Recovery enriching stage [–] 0.43Diameter of columns [m] 0.03 Stages stripping [–] 3Length of columns [m] 0.69 Recovery stripping stage [–] 0.70Feed rate [m3/h] 0.00018 Feed rate [m3/h] 0.00030Eluent rate [m3/h] 0.0085Pressure [bar] 5 Pressure [bar] 40Temperature [C�] 50 Temperature [C�] 50Solids output [kg/m3] 18.9 Solids output [kg/m3] 89

Fig. 7. McCabe–Thiele diagram for NF270 membrane cascade for xylose (lowpurity, red = stripping section, blue = enriching section, upper dashed line = equi-librium line). (For interpretation of the references to colour in this figure legend, thereader is referred to the web version of this article.)

Fig. 8. Comparison of costs for the separation of xylose from glucose with SMB andan equivalent membrane cascade with the NF270 membrane (low purity).

J. Vanneste et al. / Separation and Purification Technology 80 (2011) 600–609 607

the membrane and resin cost constitute only a small part of the to-tal cost of 0.4% and 12%, respectively. The cost of the columns (85%)and the pumps (82%) are the largest costs for the SMB and the cas-cade, respectively.

With the same SMB equipment, but with other operatingconditions, a much higher purity of 99.9% can be achieved. Thishigh purity can only be achieved by giving in on the yield, which

Fig. 9. McCabe–Thiele diagram for NF270 membrane cascade for xylose (highpurity, red = stripping section, blue = enriching section, upper dashed line = equi-librium line). (For interpretation of the references to colour in this figure legend, thereader is referred to the web version of this article.)

Fig. 10. Comparison of costs for the separation of xylose from glucose with SMBand an equivalent membrane cascade with the NF270 membrane (high purity).

608 J. Vanneste et al. / Separation and Purification Technology 80 (2011) 600–609

drops from 86% to 79%. It is impossible to get to this high puritywith a membrane cascade with only three enriching stages. If threemore stages are added a purity of 99.4% can be achieved (Fig. 9). Asthe pump cost comprised already 82% of the cost of the low puritycascade, it can be expected that with three extra stages the costwill increase significantly. From Fig. 10 it can be seen that nowboth SMB and the cascade have very similar costs. The cost ofthe SMB decreased slightly from 1921 to 1911 € per year whilethe cost for the cascade increased from 1407 to 1872 € per year.

4. Conclusion

Recently a new strategy to improve the separation performanceof membrane processes for solute–solute separations has beenadopted: the integration of membranes with moderate selectivityinto membrane cascades. This new strategy was applied here onthree different sugar separations. In order to be able to comparethe techno-economic performance of membrane cascades withthe state of the art technology, in this case SMB chromatography,a method was developed to design the membrane cascades forthe same specifications as a reference SMB installation. In thispaper the McCabe–Thiele method, previously used for the design

of gas separation membrane cascades, was adapted for membranecascades for solute–solute separations. An algorithm wasdeveloped that returns for given single-stage membrane retentionsand flux, a given feed purity and concentration and last but notleast a given productivity, yield and purity requirement from thereference SMB, the membrane cascade configuration with the low-est mixing losses.

For all studied applications a membrane cascade could be de-signed to reach the same specifications as the reference SMB. Thisis especially remarkable for the very challenging glucose–fructoseseparation where starting from 50% fructose purity a 94% fructosepurity could be reached. Due to the high number of requiredstages, for this separation the cascade cost was several timeshigher than the cost of the reference SMB. However, for the raffi-nose–sucrose and glucose–xylose separation the cascade costwas similar or lower than the cost of the SMB. Although the SMBconfigurations found in literature may not be the most profitable,the chance that the membrane and the filtration conditions ob-tained from literature are the most optimal for the studied separa-tions is also rather small. Therefore it can be concluded thatmembrane cascades are not only promising from a technical pointof view but may also be competitive relative to SMB chromatogra-phy. Moreover from the raffinose–sucrose separation and purifica-tion it could be concluded that this competitiveness increases withthe plant size. Also if the purity requirement becomes less strin-gent the competitiveness increases as could be seen from the glu-cose–xylose separation. As a result membrane cascades seem mostpromising for large scale continuous processes for producing purebut depending on the selectivity not ultrapure products. A hybridmembrane cascade SMB process could be envisaged to also coverthe ultrapure products range in a cost effective way.

Acknowledgement

IWT – Vlaanderen (Institute for the Promotion of Innovation byScience and Technology in Flanders) is acknowledged for grantingthis research.

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