Process Considerations and Economic Evaluation of 2-Step SEP for Production of Fuel Ethanol From Softwood

7/25/2019 Process Considerations and Economic Evaluation of 2-Step SEP for Production of Fuel Ethanol From Softwood

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Process Considerations and Economic Evaluation of Two-StepSteam Pretreatment for Production of Fuel Ethanol from Softwood

Anders Wingren, J ohanna So1derstro1m, Mats Galbe, and Guido Zacchi*

Depa rt ment of Chemica l Engineering, L und U n ivers it y, P . O. B ox 124, SE-221 00 L und, S w eden

To increase the overall ethanol yield from softwood, the steam pretreatment stagecan be carried out in two steps. The two-step pretreatment process was evaluatedfrom a techno-economic sta ndpoint an d compared with the one-step pretreat mentprocess. The production plants considered were designed to util ize spruce as rawma teria l a nd h a ve a capa city of 200 000 tons/yea r. The tw o-step process result ed in ahigher ethanol yield and a lower requirement for enzymes. However, the two-stepprocess is more capita l-intensive an d ha s a h igher energy requirement . The estima tedetha nol production cost w a s t he sa me, 4.13 SEK /L (55.1 /L) for both a lterna tives.For the two-step process different energy-saving options were considered, such as ahigher concentra tion of wa ter-insoluble solids in th e filter cake before the second step,and the possibility of excluding the pressure reduction between the steps. The mostoptimist ic configurat ion, with 50%wa ter-insoluble solids in t he filter cake in th e feedto the second pretrea tment step, no pressure reduction between the pretrea tment steps,

a nd 77%overa ll etha nol yield (0.25 kg E tOH /kg dry wood), result ed in a productioncost of 3.90 SE K/L (52.0 /L). This sh ows t he potent ia l for th e tw o-st ep pretr ea tm entprocess, which, however, remains to be verified in pilot trials.

Introduction

Th e e n zy m a t i c p r oce s s f or t h e p r od u ct i on of f u elethanol f rom l ignocellulosic materials has been recog-nized a s t he most promising option in terms of etha nolyield and low production cost (1-3). In this process, thepretreatment step is of vital importance because cellulosei n i t s n a t i v e f o r m c a n n o t b e h y d r o l y z e d t o a n y g r e a te xt e n t b y e n zy m e s . I n s of t w o od t h e ce ll u los e i s t h e

primary suga r source for etha nol production, and ma n-nan, the main consti tuent of the hemicelluloses, is thes e c o n d . I t h a s b e e n s h o w n t h a t c e l l u l o s e a n d m a n n a nhave dif ferent optima in terms of sugar recovery in th ep r et r e a t m e n t s t e p (4, 5). Cellulose requires a highertemperat ure th a n hemicellulose for optimal su ga r recov-ery. D uring steam pretreatment, th e sugars formed maybe degraded. Glucose, which is l iberated from the cel-lulose, is furth er degraded to 5-hydroxymethylfurfural(HMF ), levulinic acid, a nd formic acid. The hy drolysis ofhemicellulose generates pentoses, which are convertedto furfural and formic acid. These degradation products,together with l ignin degradation products and releasedorganic acids, act a s inhibitors in th e fermenta tion step(6-9) a n d i n t h e e n zy m a t i c h y d r o ly s is s t e p (10). More

severe pretreatment conditions wil l cause greater deg-rada tion of hemicellulosic sugars with loss of yield a ndpossible inhibition as a consequence (7, 9, 1 1). H owever,a rather high severi ty is required to enhance the enzy-matic digestibi l i ty of the cellulose (8, 12) . S t e a m p r e -treat ment is improved by using an acid cat alyst such asH 2S O4 or SO 2. The acid increases the recovery of hemi-cellulosic sugars and improves the enzymatic hydrolysisof the solid residue (13-17).

The dif ference in optima l pretreatment conditionsb et w e e n ce ll u los e a n d h e m ice ll u los e h a s l ed t o t h eproposal of a t wo-step pretrea tment process in w hich thefirst step is optimized for high recovery of hemicellulose,(i .e ., m a n n o s e) a n d t h e s ec on d s t e p i s op t im i z ed f orglucose recovery (14, 18, 19). The f irst step should beperformed at low severity to hydrolyze the hemicellulose.The remaining solid ma terial from the first st ep may thenbe w ashed to recover hydrolyzed hemicellulosic sugars

an d to a void further degrada tion of sugars to inhibit ingsubstances. The solid material is then treated again inthe second pretreat ment step to soften t he structure ofthe cellulose in order to increase t he enzyma tic digest-i bi li t y i n t h e s u b se q u en t p r oce s s s t e ps . Th e op t i m a lconditions for two-step steam pretreatment of softwoodha ve been investigat ed by Boussaid et a l . , Nguyen et al . ,a nd S od ers tr om et a l. (5, 12, 20-22). Impregnat ion w ithS O2 has been shown to result in higher ethanol yieldst h a n w h e n H 2S O4is used (23). Sulfur dioxide is th us th eacid catalyst used in this evaluation.

Th e p ur pos e o f t h i s s t u dy w a s t o i nv es t ig a t e t h ee con om i ca l f ea s i b i li t y o f t w o -s t e p p r et r e a t m e n t a n dcompare it to a one-step process. Since the outcome oft h e p r e t r e a t m e n t s t e p a f f e c t s t h e w h o l e p r o c e s s , i t i simporta nt not t o restr ict the evaluat ion to the pretreat-ment step only; therefore t he subsequent process st epswere also included. Process data for the two processesconsidered w as based on experimenta l w ork car r ied outa t t h e D e pa r t m e n t of C h e mi ca l E n g in ee ri ng , L u n dU n i v e r s i t y , S w e d e n (5, 23, 24). The da ta were imple-mented in t he f lowsheeting program Aspen P lus (25) t osolve the mass and energy balances in the process.

Base Case Process Description

The process evaluat ed, based on steam pretreatmenta n d s i m ul t a n e ou s s a c ch a r i f ica t i on a n d f er m e n t a t i o n

* To w hom correspondence should be a ddressed. Tel: +46 (0)-46 82 97. F a x : +4 6 (0 )4 6 4 5 2 6. E-m a il : g u ido.z a c ch [email protected].

1421Biotechnol. Prog. 2004, 20, 14211429

10.1021/bp049931v CCC: $27.50 2004 American Chemical Society and American Institute of Chemical EngineersPublished on Web 08/26/2004


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(S S F ), i s i ll us t r a t e d i n F i gu r e 1 . Th e r a w m a t e r ia lconsists of f resh spruce w ith a wa ter content of 50%.Eight area s w ithin the production plan t a re considered:feedstock ha ndling, pretreat ment, S SF, dist i l lat ion, de-w a t e ri ng , e va p or a t i on , s t ea m g en er a t i on , a n d ot h e rfacilities. The la tt er consists of off-site fa cilities such a sstorage fa cil i t ies for chemicals and products a nd a dmin-istrat ion buildings.

Pretreatment. A flowsh eet for t he pretreat ment st epsca n b e s e en i n F i gu r e 2 . F r es h ch ip s o f s pr u ce a r et r a n s p or t e d f r om t h e r a w m a t e r i a l h a n d l i n g a r e a t o t h ep re t re a t m en t s t ep w h e re t h e y a r e i m pr eg n a t e d w i t hs u lf u r d i ox id e b ef or e b e in g f ed t o t h e p r et r e a t m e n treactor . High-pressure stea m, a round 22 bar , is used totreat the w ood chips a t t he desired temperatur e for the

residence time required. The steam consumption in thepretreatment units was set 10%higher than for a systemw o r ki n g a d i a b a t i c a l ly t o a cc ou n t f or h e a t l os s es a n dsteam leakage in the steam distr ibution system.

After pretrea tment t he mat erial is flashed in tw o steps.A f t e r t h e f i r s t f l a s h i n g t o 4 b a r t h e v a p o r f o r m e d i scondensed a t 138 C in hea t exchanger 1 (HX1) a nd th enfurther f lashed to a tmospheric pressure a nd condensedat 95 C (in HX2). In the third heat exchanger (HX3) thevapor at 1 bar, from the second flashing of the pretrea tedma teria l, is condensed a t 95 C. The purpose of reducingthe pressure in t wo steps is to recover some of the la tentheat at a higher temperat ure. The second pretreat ments t ep i n t h e t w o -s t ep p roce ss i s s im il a r t o t h e f ir s t .Dewa tering and w ash ing steps are included after the first

pretreat ment step utilizing recycled evaporation conden-s a t e s a v a i l a b le a t 6 0 C a s w a s h i n g w a t e r . Th e con t e n tof w a t e r -i n s ol u bl e s ol id s (WI S ) i n t h e f il t e r ca k e i sa s s u m e d t o b e 3 0% i n t h e b a s e c a s e p r oce s s, b u t t h i spara meter wa s varied to study the impact on the ethanolproduction cost . In t he simulat ions i t w as assumed t ha tf i l trat ion is carr ied out in a counter-current fashion int w o s t a g e s , a n d t h e r e l a t i o n b e t w e e n t h e a d d i t i o n o fw a s h i n g w a t e r a n d w a s h i n g e f f i c i e n c y w a s b a s e d o ns t u d i es of t h e r e m ov a l of l ig n i n f r o m p u lp (26). Thea m o u n t o f w a s h i n g w a t e r w a s a d ju s t e d t o r e m ov e 95 %o f t h e s u g a r s , a n d s u g a r s f e d t o t h e s e c o n d s t e p w e r eassumed to be una f fected, i .e. , not degraded t o byprod-ucts. The w hole f i ltra te from the wa shing step t ogetherwith the slurry leaving the second step is tra nsferred to

t h e S S F s t e p.An a lterna tive to th e tw o-step process described a boveis to exclude the fla shing aft er the first pretreat ment st epa n d t o d ew a t e r a n d w a s h t h e m a t e r ia l u n d e r p r es s u r e.The f i l trate in this process configuration is f lashed toatmospheric pressure, and the latent heat in the vapori s r e cov er e d a t 9 5 C (H X 8 i n F i g u re 2 ). B e c a u s e t h ep r e s s u r e , a n d t h u s t h e t e m p e r a t u r e , o f t h e s l u r r y i smaintained, the amount of steam needed in the seconds t e p i s d e cr e a s e d . Th e a d d i t ion of w a s h i n g w a t e r a l s or e su l t s i n a d e cr e a s e i n t h e t e m p er a t u r e of t h e s l ur r ybeing fed to th e second step. It ha s been shown, however,tha t suga rs not w ashed out, e.g. , those fed to t he seconds t ep , a r e d eg r a d ed t o a l ow e r e x t en t t h a n m i gh t b eexpected. I n a stud y by Soderst rom et a l. (27) tw o-step

Figure 1. Simplif ied f lowsheet for th e process considered in t his s tudy .

Figure 2. One-step and two-step pretreatment.

1422 Biotechnol. Prog., 2004, Vol. 20, No. 5


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pretreatment with and without dewatering and washingb et w e e n t h e p r et r e a t m e n t s t e ps w e r e c om p a r e d u s i n gS S F f or t h e e va l u a t i on of t h e ov er a l l e t h a n ol y i el d.

Although the ma nnose recovery af ter pretreat ment w ashigher with dewatering and washing between the steps,t h e o v e r a l l e t h a n o l y i e l d w a s a b o u t t h e s a m e f o r b o t hcases. Thus, a case in which the washing was excludedwhile assuming t he overal l ethan ol yield to be the sa mea s i n t h e t w o -s t e p b a s e ca s e w a s a l s o e va l u a t e d i n t h i ss t u d y .

P retreatm ent da ta for the one-step base case is basedon experimental da ta presented by G albe et a l . (24) Thep r et r e a t m e n t t e m pe r a t u r e w a s 2 15 C , r e s id e n ce t i m ewa s 5 min, and S O2wa s added a t 2%of the wa ter contentin the wood. These conditions correspond well to theop t im u m i n t e rm s of ov er a l l s u g a r y i el d, a s w e ll a sethanol yield, as determined by Stenberg et al . (18).

Da ta for tw o-step pretrea tment originat es from studies

car ried out by S od erst rom et a l. (5, 23). Various condi-tions were studied in terms of overal l ethan ol yield, andthe optimal conditions were found to correspond to a firststep at 190 C for 2 min an d a second step at 210 C for5 m i n . B ot h p r et r e a t m e n t s t e ps w e r e p r e ce d ed b y i m -pregna tion w ith 2% sulfur dioxide based on th e l iquidcontent in t he w ood.

The pretreatment was modeled as a reactor workingwit h fixed fra ctiona l conversion of the components in t hewood. The reactions considered were hydrolyses of hex-osans and pentosans to monomers, their degradation toH M F a n d f u r fu r a l , a c et i c a ci d p r od u ct i on f r om a c e t y lgroups in the wood, a nd lignin degra da tion. Two lumpedcom p on e n t s , on e v ol a t i l e a n d on e n on v ol a t i l e, w e r edefined to account for further degra da tion products such

as formic acid, levulinic acid, and unknown products. Thecomposition of the raw ma teria l and t he yields in the tw op r e t r e a t m e n t a l t e r n a t i v e s a r e g i v e n i n F i g u r e 3 . T h e

main difference between the two process configurationsis the higher fra ctiona l conversion of mann a n to ma nnosei n t w o -s t e p p r et r e a t m e n t co mp a r e d t o t h e o ne -s t e pprocess. Furthermore, a larger proportion of the cellulosei s con v er t e d t o g l uc os e i n t h e t w o -s t e p p r e t r ea t m e n tprocess th a n in t he one-step process. However, in t he tw o-step process slight ly more of the cellulose is degrad ed tobyproducts, w hich cannot be used for eth a nol production.Of t he further degraded components 85% is volati les,most of which end up in the f lash steam together with50%of the degra ded lignin. For t he tw o-step process th eamount of volatile lignin was set to 25%of the degradedlignin, expressed as weight, which is the same amountas for the one-step process.

Simultaneous Saccharification and Fermenta-

tion. I t w a s a s s u m ed t h a t S S F w a s p e r fo rm e d in a f ed -b a t c h m o d e a n d p r o c e s s d a t a w a s b a s e d o n t h e s a m eexperimental series as for the one-step pretreatment (24).In these experiments, commercial ba kers yea st, Saccha-romyces cerevisiae, wa s a dded t o a concentr a tion of 5 g/L.In a commercial plant a more real ist ic option is to uses om e o f t h e s u g a r s a v a i l a b l e i n t h e h y d r o ly s a t e f or i n -house yeast production. In this study it wa s assumed th atyeast must be produced in-house to a concentration of 2g /L i n S S F u s in g f er m e n t a b l e s u g a r s a v a i la b l e i n t h ehydrolysate from the pretreat ment st ep. The a mount ofsugar needed w il l thus be dependent on the volumetricf low r a t e t o t h e S S F s t e p. A s a r e s ul t o f t h i s i n -h ou s eyeast production the calculated ethanol yield will be lowerthan the yield obtained in the experiments. At the end

Figure 3. C omposition of raw ma terial a nd yields af ter pretreat ment, expressed as g/100 g dry ra w m at erial . For the tw o-step caseyields are given as overall yields from both the f irst and the second step.

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of SS F the tota l amount of f reshwat er added is the sameas i f the process had been carr ied out in conventionalba tch mode w ith a n init ial loa d of WIS of 8.4%. The loadof cellula ses is 15 FP U /g WIS, a nd 95%of th e glucan ishydrolyzed to glucose. Ninety-two percent of the ferment-a b l e s u g a r s ( g l u c o s e a n d m a n n o s e ) t h a t r e m a i n a f t e ryeast ha s been produced is fermented to ethan ol, and t herema inder (8%) is converted to bypr oducts, e.g., g lycerol,o r r e m a i n u n f e r m e n t e d . I n t h e e x p e r i m e n t s t h a t t h i sstudy is based on, commercial ba kers yeast w as used a s

t h e f e r m en t a t i on or g a n i s m . B e ca u s e t h i s y e a s t ca n n o tferment pentoses to ethanol , i t was assumed that thesesugars will remain unfermented. The total reaction timein the S SF step is 72 h, a nd t he tota l cycle time, includingdraining and cleaning, is assumed to be 84 h.

Fed-batch SSF studies, similar to those described abovefor the one-step process, have not been carried out witht w o - s t e p p r e t r e a t e d m a t e r i a l . H o w e v e r , t o b e a b l e t ocompare the processes i t was assumed that the resultsin terms of yields in the S SF step would be the same forboth processes. This assumption is justified by the factt h a t t h e y i el d s ob t a i n e d in b a t ch S S F , w i t h a s u b st r a t eload of 5%, of mat erials pretreated in one step and tw os t e ps w e r e b o t h a r o u n d 9 0% b a s e d o n t h e a m o u n t s o fg l uca n , g l uco se , a n d m a n n o s e f ed t o t h e S S F s t e p (23,24).

Distillation and E vaporation. Th e d ow n s t r e a mprocessing begins with a distillat ion step, which consistso f t w o s e p a r a t e s t r i p p e r s a n d a r e c t i f i e r o p e r a t i n g a tdif ferent pressures. The overhead vapor from the f irststripper is used as heating medium in the reboiler of thesecond stripper, an d th e overhea d va por from th e secondstripper is used as heating medium in the rectifier. Thep rod u ct i s e t ha n o l a t 94 w t %. N o f u r t h er e t h a n olpurification was considered in the analysis. The stillagefrom the strippers is filtered, and the solid fraction is sentto the drying st ep. The l iquid fraction, with a dry ma tt ercont ent of around 3.5%, is concentra ted t o a dry m a tt ercontent of 50%in a multiple-effect evaporator consisting

of f ive stages with a forward feed arrangement.Drying, Steam Generation, and Pellet Produc-

tion. The solid residue from the f i l tra t ion step a nd t heevaporation residue are mixed and dried in a stea m dryerto a d ry ma tt er content of 85%. Pa rt of this dr ied solid isincinerated to produce the steam needed in the pretreat-m e n t , t h e d i s t il la t i o n , t h e e va p or a t i o n , a n d t h e d r y i n gsteps. The excess material is pelletized and sold as a solidfuel.

Methods

The methodology used in t his w ork is similar to th atdescribed previously (28); however, a brief description,including the updates to the model , is provided here.

Aspen P lus from Aspen Technology (25) w a s u s e d t os i m ul a t e t h e p r oce s s. R i g or ou s m a t e r i a l a n d e n er g ybalance calculations using detai led equipment modelswere carr ied out to determine f low rates, composit ion,and energy f low for a l l streams in the process. Sizing oft h e e q ui pm en t w a s ca r r i ed ou t u t il iz in g t h e I ca r u sProcess Evaluator (IPE) (29) a nd by using r ules of thumb(30). E q u i pm en t cos t w a s e st i m a t ed u s in g t h e I P E ,various reports, and vendor q uotations. The cost of thep r e t r e a t m e n t r e a c t o r s w a s b a s e d o n a q u o t a t i o n f r o mSt a ke Technology (31). To evalua te t he consistency of theequipment cost provided by the IP E, vendor quota tionswere obta ined from a Sw edish engineering company . Theprocess equipment compared was columns, pumps, heatexchangers, and vessels including fermentors. The cost

e st i m a t ed w i t h t h e I P E d ev ia t e d b y (20% from the

vendor costs, w ith t he largest deviat ion for the sma lleste q u i p m e n t . F o r t h i s k i n d o f a c o m p a r a t i v e s t u d y t h i sdeviat ion is a ccepta ble. For quoted equipment, i .e. , thepretreat ment rea ctor a nd pellet ma chine, and for equip-ment costs obtained from the l i terature, i .e. , the dryera n d s t ea m b oi le r, t h e s ix t h -t e n t hs r u le w a s u s ed t ocalculate the equipment cost , C2, a t a c a p a c i t y , A 2, fromdata at a specif ied capacity, A 1, with a cost C1:

The IP E w a s also used to estima te the cost of auxiliaryequipment, such a s piping a nd instrum entat ion, to yieldthe direct cost . Indirect capital , w hich consists of costsnot directly relat ed to the process equipment its elf , suchas engineering, f reight, star t-up cost , and contingencycosts, wa s also estimated w ith the IP E. The f ixed capita linvestment consists of th e sum of the direct a nd indirectcos t s . W or k in g ca p i t a l w a s e s t im a t e d a s s u g g es t e d b yPeters and Timmerhaus (32).

The annual cost of the f ixed capital was obtained bymultiplying the f ixed capital investment by an annual-ization factor of 0.104, corresponding to an interest rateof 6%a nd 15 years pay-off t ime. Zero salva ge value w asassumed. Ra w mat erial cost , 90 SEK /MWh, w as basedon the a ctual cost of forest residue in Sw eden, w hereasthe assumed income from the solid fuel coproduct, 140

SE K/MWh, is sl ightly lower tha n the current price forpellets an d briquettes. The cost of enzymes, 19 SE K/million FPU, was based on an estimate for producing thee n zy m e s a t t h e p la n t (33). All costs are summarized inTable 1.

Th e p la n t w a s a s s u m ed t o b e a g r a s s -r oot s p la n tdesigned t o process 200 000 tons of dry ra w ma terialyea rly. The on-line time w a s set to 8000 hours per yea r.The product wa s a ssumed to be 94%etha nol, but a ll costsan d energy consumptions reported a re based on a puree t h a n o l p r o d u c t . C o s t s r e p o r t e d i n U S $ ( o r U S ) a r ebased on a n excha nge ra te of 7.5 SEK /US $. The a im oft h e s t u d y w a s t o c om p a r e d i f fe r en t p r et r e a t m e n t con -figura tions. The calculat ed costs should th erefore not ber e g a r d e d a s a b s o l u t e , a n d c a r e m u s t b e t a k e n w h e n

Table 1. Costs Used in the Evaluation.

t ype cost un it

r a w m a t e r ia lw ood 0.42 S E K /kg D M

chemicalssulfur dioxide 1.5 S E K /kglim e 0.70 S E K /kgdefoa m er 20.0 S E K /kgsodium h y dr oxide 2.00 S E K /kg(NH 4)2H P O4 1.50 S E K /k g

MgSO4 4.41 S E K /k gen zy m es 19.0 S E K /106 F P Ubyproduct income

solid fuel 0.79 S E K /kg D Mca r bon dioxide 0.03 S E K /kg

utilitieselect r icit y 250 S E K /MWhcooling w a t er 0.14 S E K /m 3

pr ocess w a t er 1.40 S E K /m 3

other costsl a bor (per em pl oy ee) 500 000 S E K /y ea rin sur a n ce 1 %of fixed ca pit a l

investmentm a in t en a n ce 2 %of fixed ca pit a l

investment

C2 ) C1(A 2

A 1)

0.6



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comparing the costs obtained in this study with resultsfrom similar studies based on other a ssumptions.

Results and Discussion

Base Cases. Strea m da ta f rom the simulations for thetw o bas e cases is shown in Ta ble 2. The yield of WIS (kgWIS /kg ra w ma teria l) a fter pretrea tm ent is 0.61 a nd 0.56for the one-step and two-step process, respectively. Forboth process configurations the concentra tion of WISa f t e r f l a s h i n g i s h i g h e r i n t h e s i m u l a t i o n s t h a n i n t h ee x p e r i m e n t s , a n d t h e d i f f e r e n c e i s m a i n l y d u e t o t h eh ig h er h ea t l os s i n t h e s m a l l l a b or a t o ry e q ui pm en tcompared to that in the simulated full-scale, continuouslyworking rea ctor. A high hea t loss results in condensa tionof the f lash ed vapor and thus di lution of the slurry.

For the one-step base case the energy requirement inthe pretreatment sta ge is 14.1 MW, w hich correspondsto 0.73 kg st eam /kg d ry wood, see Ta ble 3. The energ yduty associated with the condensation of f lash vapor is

in tota l 6.9 MW, wh ich results in a net energy require-m e n t f or t h e p r et r e a t m e n t s t e p o f 7 . 2 M W. Th e t o t a le n er g y r e q u ir e m en t f or t h e w h o le p la n t i s 3 8. 3 M W.Sixty -one percent of the solid residue from the dr yer mus tthus be incinerated to produce the required amount ofs t e a m .

Two-step pretreatment requires 26.7 MW or 1.37 kgstea m/kg dry w ood, a nd t he recovera ble energy from thef lash vapor is 14.8 MW, which results in a net energyr e q u ir e m en t of 1 2. 0 M W. Th e t o t a l p r oc es s e n er g yr e q u ir e m en t i s s l ig h t l y h i g h er t h a n f or t h e on e -s t e pprocess; 40.6 MW and 64%of the solid residue must beincinera ted. The difference betw een th e tw o processes interms of tota l process energy requirement, 2.3 MW, islower than would be expected from the difference in the

pretreatment step only, which is 4.8 MW. However, theenergy requirement in the downstream processing stepsis lower for the two-step process than for the one-stepprocess as a result of less f reshwater being needed inSS F. This is a result of the lower y ield of WIS in the t wo-step process than in the one-step process.

Ut i l izat ion of t he secondary steam requires low-tem-pera ture process steps for t he steam to ha ve a value. I fthis is not possible, the energy in t he vapor formed mustbe removed by condensa tion using cooling wa ter, a nd th enet energy requirement for pretreat ment w il l be higherthan that calculated here. In the processes consideredin this study, the downstream processing steps requiresubstantial amounts of steam. For instance, the f lashedvapor can be used to preheat the feed to the dist i l lat ionstep or in the reboiler of the stripper.

The flow of etha nol after S SF in th e one-step ba se caseis 5860 kg/h (0.23 kg Et OH /kg d ry w ood), wh ich is 71.8%of the t heoretica l maximum tha t could be produced fromthe fermenta ble hexoses (glucose and ma nnose) a vailablein the raw material . I f no yeast had been produced, thee t h a n o l y i e l d w o u l d h a v e b e e n t h e s a m e a s t h e y i e l dobtained in the experimental study, i .e. , 76%(24). Theetha nol production in th e tw o-st ep bas e case is 6090 kg/h(0.24 kg Et OH/kg dr y w ood), corres ponding t o a y ield of74.6% with yeast production an d 78.1% with out yeastproduction. Tw o-step pretrea tm ent followed by ba tch S SFa t 5% WI S h a s r es u lt e d i n a n e xp er im en t a l ov er a l le t h a n o l y ie ld o f 8 1% b a s e d o n t h e g l u ca n a n d m a n n a na v a i la b l e in t h e r a w m a t e r ia l (23).

I t w a s a s s um ed t h a t t h e s u ga r s f ed t o t h e s econ dpretreatment reactor were not degraded. In the base casetwo-step process 95%of the sugars was removed beforethe second pretreatment step. I f 20%of the sugars fedto the second pretreatment reactor were degraded, the

overall ethanol yield would be reduced by only 0.2%agepoints, w hich justifies the a ssumption of no degra da tion.Ta b l e 4 s u m m a r i z es t h e c os t s a s s o ci a t e d w i t h t h e t w obase cases evalua ted, showing a n overall ethanol produc-ti on cost of 4.13 SE K/L (55.1 /L) for bot h pr ocesses. Thema jor adva nta ges of the tw o-step process a re i ts highere t h a n o l y i e ld a n d i t s l ow e r e n zy m e r e q u ir e m en t . Th ea mount of enzymes a dded is based on the a mount of WISfed to SS F. As more materia l is hydrolyzed in tw o-steppretreatment, less WIS is present in SSF and thus lessenzyme solution is needed. The lower amount of WISa fter pretreat ment in tw o steps also reduces the need forfreshwat er to the SSF step, thereby reducing the capitalcost and energy consumption in downstream processing.The cost of the enzymes was in this study estimated to

Table 2. Composition of Streams in the Base Cases of One- and Two-Step Pretreatmenta

on e-st ep t w o-st ep

st r ea m 1 2 3 4 5 2 3 4 5 6 7 8 9 10 11

liquid phasew a t er 25.0 18.2 7.6 3.6 31.1 14.9 4.9 3.6 30.2 38.7 47.8 19.5 9.8 4.7 90.5h exoses 4.8 4.6 0.2 4.4 6.1pen t oses 0.0 0.0 1.4 0.0 0.0 1.7 0.1 1.6 0.0 0.0 1.8ot h er s 0.9 1.4 0 .4 1 .4 0.8 0 .3 1 .0 0.1 1.0 0.3 0 .1 1.7

solid phase

h exosa n s 15.0 8.6 10.0 10.0 7.8pen t osa n s 1.8 0.1 0.2 0.2lign in 7.0 6.6 6.6 6.6 6.2a ce ty l g r ou ps 0. 2

total flow 50.0 18.2 8.9 4.0 53.9 14.9 5.6 3.9 54.3 55.9 54.7 19.5 10.1 4.7 114.0

t em p ( C ) 25 217 144 100 100 217 144 100 100 70 82 217 144 100 91pr essur e (ba r ) 1 22 4 1 1 22 4 1 1 1 1 22 4 1 1

a Flow ra tes in ton/h. St ream numbers refer to Figure 2.

Table 3. Energy Duties (MW) in the Pretreatment Step

on e-s t ep t w o -s t ep p r es s ur e (b a r )

s t e a m 1a 14.1 11.6 22s t e a m 2a 15.2 22H X1 -4.3 -2.8 4H X2 + H X3 -2.6 -2.5 1

H X4 + H X8 0.0H X5 -5.9 4H X6 + H X7 -3.5 1net energy demand

in pretreatment7.2 12.0

steam requirement(kg/kgdry wood)

0.73 1.37

a Ca lc u la t e d a s t h e h e a t n e ed e d t o p r odu ce t h e s t e a m f r om aboiler fed water a t 25 C.



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be 0.6 SE K/L E tOH (8.0 /L E tOH ) for t he one-st ep ba secase. Himmel et a l . (34) have a lso estima ted t he cost forenzymes produced from hydr olysa te slurry. F or a plantcapacity of 2000 dry tons per d ay of feedstock t he costwa s 9 /L E tOH. Na tional renewa ble energy laborat ory(NRE L) expects a cost of 2.6 /L E tOH (35). H owever, int h e s e t w o s t u d i e s a l o w e r d o s a g e o f e n z y m e w a s u s e dt h a n i n t h e p r es e n t s t u d y .

The f ixed capita l investment w as estimated to be 847and 893 million Swedish kronor (MSEK) for the one-stepa nd t he tw o-step processes, respectively. The ma in rea sonfor the t wo-step process being more expensive is t he extra

pretreatment r eactor and the ad ditiona l fla sh vessels andthe heat exchan gers. The dewat ering and w ash ing step,however, w as not included in t he economic evaluat ion.The higher cost of pretreatment is somewhat of fset byt h e l ow e r ca p i t a l cos t of d ow n s t r e a m p r oce s si n g . I naddit ion t o the higher capital cost , the tw o-step processsuffers from a higher consumption of sulfur dioxide an da lower income from the solid fuel coproduct. The la tt eris due to t he higher energy consumption.

Process Alternatives for the T wo-Step Process.In the ba se case tw o-step process t he slurry is f lashed,dewatered to 30%WIS, and washed with recycled con-densate from the evaporation unit to recover 95%of thesugars. For these conditions the net energy duty in the

pretreatment step is 40%higher in the t wo-step base casetha n in th e one-step base case. In a n a t tempt t o reducethe energy consumption in t he tw o-step case, th e depen-dency of the concentration on WIS in the filter cake (a)an d the impregna tion tempera ture (b) on the energy dutywa s st udied. In ad dit ion, t he consequence of excludingthe f lashing (c) and the washing (d) between the stepsare discussed.

( a ) H i g h e r C on c en t r a t i o n o f W I S i n t h e F i l t er

C a k e . B y increasing t he fraction of WIS in the filter cakeit is possible to reduce the energy consumpt ion, since lessl iq u id n ee ds t o b e h e a t e d i n t h e s econ d s t ep . M o reefficient dewa tering a lso reduces t he a mount of cat alystneeded in t he impregnation preceding t he second step,

a s t h e a m o u n t i s b a s e d o n t h e w a t e r c o n t e n t o f t h em a t er ia l . F i g ur e 4 s how s t h e n e t e ne rg y d ut y a s afunction of concentra tion of WIS in t he filter cake for th etw o-step pretreatm ent (Case a 1). If th e concentra tion canbe increas ed, from th e ba se ca se va lue of 30% to 50%,the net energy duty will be reduced from 12.0 to 8.9 MW,a reduction of 26%. A WIS concentra tion of 70% wouldr ed u ce t h e n et d u t y b y 37% t o 7 .6 M W. H ow e ve r,d e w a t e r i n g a n d f il t r a t i on of t h e m a t e r ia l f r om t h e f i r s tpretrea tment step ha ve not been technical ly optimized,a WIS concentra tion of 37%being rea ched on lab scale(21). A higher concentration is expected for an optimized,large-scale system. Fifty percent WIS was therefore takenas the currently most real ist ic value.

( b ) I m p r eg n a t i o n T em p er a t u r e . After the first stept h e s l u r r y h a s a t e m pe r a t u r e of 1 00 C , b u t a s a r e s ul tof t h e a d d it i on of w a s h in g w a t e r, t h e t em per a t u redecreases. A higher concentration of WIS in the f i l tercake results in a higher temperature of the feed to thesecond reactor due to the lower addition of cold washingwater . A WIS concentration of 50%yields a feed of 85 C, w hereas a WIS concentra tion of 30%corresponds t o

7 0 C . Th e t e m pe r a t u r e , h o w e ve r , a f fe ct s t h e c a t a l y s timpregna tion of the ma terial . For a l iquid cata lyst suchas sulfuric acid impregnat ion w il l probably be better a ta high temperat ure as a result of a higher dif fusion ra te.However, for a gaseous cata lyst , such a s sulfur dioxide,the t empera ture ha s to be low for the ga s to dissolve inthe liquid. From a process point of view this means thatt h e s l u r r y , a f t er b ei n g w a s h e d , m a y h a v e t o b e c oo le dfurther before impregnation can take place. Thus, thee ff ect of h a v i n g t o coo l t h e s l ur r y b e fo re t h e s ec on dimpregna tion sta ge w as investigat ed. Two WIS concen-trations in the f i l ter cake, 30% and 50% (Case b1 andCase b2), were studied. The results are shown in Figure5. Th e w or s t ca s e i s w h en t h e f il t er ca k e ca n n ot b edewatered to more than 30%WIS and has to be cooled

to 20 C before impregnation. For these conditions thenet energy dut y increases to a lmost 16 MW.( c ) N o F l a sh i n g a f t er t h e F i r st S t e p . An a lternative

to t he process configura tions discussed a bove is to excludet h e f la s h i n g a f t e r t h e f ir s t p r et r e a t m e n t s t e p. Th e d e -w a t e r in g , w a s h i n g , a n d i m pr e gn a t i o n s t e p s a r e ca r r i e dout under pressure, after which the slurry is fed directlyt o t h e s e con d p r e t r ea t m e n t s t e p. I t w a s a s s u m ed , a s i nt h e ca s e s d i sc us s ed a b o ve , t h a t r e cy c le d e va p or a t i o ncon d e n sa t e s a t 6 0 C c a n b e u s e d a s w a s h i n g w a t e r f or95%suga r recovery. The option of heating the wa shingstrea m to 180 C (HX4) wa s also studied. Figure 4 showst h e r es u lt s w h e re t h e con ce nt r a t i on of WI S i s a l s oincluded as a parameter. The difference, in terms of netenergy deman d, is minor betw een flash ing to 1 bar (Ca se

Table 4. Costs in the Base Case Processes

one-st ep t w o-st ep

y e arl y c ost c ost/L E tO H y e arl y c ost c ost/L E tO H

MS EK M$ S EK cen ts MS EK M$ S EK cent s

r aw m at l 83.0 11.1 1.41 18.8 83.0 11.1 1.36 18.1ch em ica ls 50.4 6.7 0.86 1 1.4 51.5 6.9 0.84 1 1.2byproducts -31.3 -4.2 -0.53 -7.1 -29.5 -3.9 -0.48 -6.4ut ilit ies 11.5 1.5 0.20 2.6 12.1 1.6 0.20 2.6ot her cost s 37.9 5.1 0.64 8.6 39.3 5.2 0.64 8.6ca pit a l 91.5 12.2 1.55 20.7 96.4 12.9 1.58 21.0

total 243 32.4 4.13 55.1 253 33.7 4.13 55.1

Figure 4. N e t e n er g y d u t y i n t h e p r et r e a t m en t s t ep a s afunction of WIS for various alternatives. Dashed lines representt h e t wo-s t e p b as e c as e an d t h e s ol id h or iz on t a l l i n e t h e on e -

s t ep b a s e c a s e. C a s e a 1 : f la s h i n g t o 1 b a r a f t e r f i rs t s t e p ,wa s h i n g wat e r a t 60 C , 95% r e cove r y of s u g ar s . C as e c 1: n oflashing af ter f irst s tep, washing water at 60 C, 95%recoveryof su g ar s . C a s e c2: n o f las h i n g a f t e r f ir s t s t e p , was h i n g wa t e rat 180 C , 95% r e cove r y of s u gar s . C a s e d 1: f las h i n g t o 1 b a r ,n o wa s h i n g . C as e d 2: n o f las h i n g a f t e r f ir s t s t e p , n o was h i n g .



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a 1 ) a n d w i t h ou t f la s h in g a n d w a s h in g w a t e r w i t h atemperature of 60 C (Case c1). However, if the washingwater is heated to 180 C (Case c2), close to the temper-ature of the slurry from the f irst step, the recovery ofe n er g y f r om t h e f il t r a t e i n cr e a s es , r e s u lt i n g i n a n e tenergy reduction.

( d ) N o W a s h i n g o f t h e F i l t e r C a k e . The differentoptions discussed above assume t ha t 95%of the suga rsreleased from the first pretreatment step are washed outof the filter cake before the second step. The option ofexcluding the washing step wa s studied with a nd with outfl a sh in g a f t er t h e f ir st s t ep, C a s e d 1 a n d C a s e d 2,respectively. Although no wa shing wa ter is ad ded, sugarsare recovered as a result of the dewatering step alone.F or i ns t a n ce , a t 50% WI S i n t h e f il t er ca k e a n d n of lashing, 65% of the sugars is removed from the f i ltercake. These two cases are also i l lustrated in Figure 4.When no fla shing is carried out aft er the first step, a verylow net energy consumption is a chieved, even a t lowconcentra tions of WIS in the f i lter cake. At 40% WIS,the net duty is even lower tha n for the one-step base cas e(


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highest experimental overall ethanol yield achieved in atwo-step process was 81%(23). This indica tes a m a nna n-to-ma nnose conversion of 85% a nd 100% hydr olysis ofcellulose according to Figure 7, represented by 77% int h e s i m ul a t i on s . Th e se a r e r e a l is t i c a s s u m p t i on s a s asuccessful SS F experiment provides an etha nol yieldabove 90%, indicat ing the complete hydrolysis of cel-lulose. Figure 8 shows t he etha nol production cost as afunction of the et ha nol yield for the t w o-step process. Theenergy consumption was assumed to be the same as inthe tw o-step base case. At t he highest yield at t ained inexperimenta l work, 77.0%, the lowest eth a nol productioncost is 4.04 S E K/L.

Summary of the Different Cases. F i gu r e 9 s u m -mar izes some of the cases discussed in this study. The

tw o-step configura tion tha t results in t he highest etha nolproduction cost, 4.23 SEK /L, is t ha t in w hich the slurr yf rom t h e f ir s t s t ep h a s t o b e c ool ed t o 20 C b ef or eimpregnat ion. The lowest cost w hen 95%of the sugars

is washed out before the second pretreatment step, 4.00S E K /L , i s ob t a i n e d w h e n p r et r e a t m e n t i s con f ig u r edaccording to Case c2 in Figure 4, with 50%WIS in thefilter cake (C95%in Figure 9).

A best case w as identi fied in which the m ost favor-a b l e co nf ig u r a t i on a n d d a t a w e r e i n cl u de d : a y i el d o f77.0% (80.5% wit hout yea st production) correspondingto a mannose recovery of 85% and 100% hydrolysis ofcellulose, an energy consumption in the pretreatmentu n it cor r es pon d in g t o t h e ca s e w h e re f la s h i ng w a s

excluded, no washing water added (Case d2), and 50%WIS in the filter cake. Und er these conditions the etha nolproduction cost w a s estima ted t o be 3.90 SE K/L, w hichis 5.6%lower t ha n t he 4.13 SE K/L in t he one-step ba secase process.

Conclusions

The two-step process results in an ethanol yield highertha n tha t of the one-step process, 74.6% compared to71.8%. Another adva nta ge is the lower requirement ofe n z y m e s a n d w a t e r i n t h e S S F s t e p . M a j o r d r a w b a c k sa r e , h o w e v e r , t h e h i g h e r c a p i t a l c o s t a n d t h e h i g h e renergy consumption. The overal l etha nol cost wa s esti-ma ted to be 4.13 SE K/L (55.1 /L) for both th e one-st ep

and the two-step base case. The lowest cost estimatedfor t he tw o-st ep process, 3.90 S EK /L (52.0 /L), req uiresa high ethanol yield, high concentration of WIS in thef i lter cake betw een th e steps, and tha t t he sugars beingfed to the second step a re not degra ded. The higher y ieldha s been demonstra ted experimenta lly, but t he tw o otherassumptions st i l l need to be verif ied.

An i m por t a n t p a r a m et e r i n t h i s e v a l ua t i on i s t h eoveral l etha nol yield, which ha s a significant impact onth e etha nol production cost. Therefore, it is of import a nceto follow the progress ma de in both t he one-step a nd t wo-step process. The difference in yield betw een the t wo ba secases, 2.8%, can serve a s a rough guide w hen comparingfuture experiments.

Another a spect of importa nce is the higher complexity

of a two-step pretreatment process compared to a one-step process. Whether it is possible to operate the two-step process in the w ay suggested in this paper must betechnically verified on a la rger sca le. Hopefully, some of

Figure8. E tha nol production cost as a function of ethan ol yieldfor the tw o-step process. Da shed lines indicat e the t wo-step basecase, and the solid l ine represents the highest yield achieved.

Figure 9. Sum ma ry of costs . One-step BC : one-step base case, 71.8% etha nol yield. Two-step BC : tw o-step base case, wa shingwa ter a t 60 C , 74.6%etha nol yield. Two-step worst case: s lurry from the f irst s tep cooled to 20 C, 30%WIS in f i l ter cake, and74.6%etha nol yield. C95%: no f lashing a f ter f irst s tep, wa shing wa ter a t 180 C, 50%WIS in f i l ter cake, an d 74.6%etha nol yield.optimistic case: no f lashing a f ter f irst s tep, no wa shing, 50%WIS in f i l ter cake, 77.0%etha nol yield.



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the questions raised in this paper can be answ ered w henthe new pilot plant that is being constructed in the northof S w e d en i s i n op er a t i on (36). I t w i ll op er a t e i n acontinuous mode (2 ton dry ra w ma terial per da y) withpretreatment in one or two steps. The dewatering andwashing steps are designed to work under pressure, assuggested in this paper .

Acknowledgment

The Swedish Energy Agency is gratefully acknowl-edged for its financial support.

References and Notes

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Accepted for publication J uly 16, 2004.

BP049931V


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