An economic evaluation of the fermentative production of lactic acid

An economic evaluation of the fermentative production of lactic acidfrom wheat ¯our

Christina �Akerberg, Guido Zacchi *

Department of Chemical Engineering 1, Lund University, PO Box 124, SE-221 00, Lund, Sweden

Received 19 December 1999; received in revised form 2 March 2000; accepted 13 March 2000

Abstract

A process for the fermentative production of lactic acid from whole-wheat ¯our consisting of starch and bran containing nu-

trients is presented and an economical evaluation of the lactic acid production cost performed. Bottlenecks were identi®ed and

alternative processes were evaluated and compared. The costs of raw material, the sodium hydroxide in the fermentation step, and

the conversion of lactate to lactic acid using electrodialysis were found to contribute considerably to the total production cost.

Performing the fermentation step as a batchwise step was economically better than continuous fermentation. The lactic acid pro-

duction cost can be reduced by lowering the pH and/or by recycling the sodium hydroxide produced by electrodialysis to the fer-

mentor. Using higher wheat ¯our concentrations reduced the lactic acid production cost and numerical optimisation of the process,

with respect to the wheat ¯our concentration, showed that the optimal concentration corresponded to 116 g glucose/l, which resulted

in a production cost of 0.833 US$/kg product. A Monte Carlo simulation of the total production cost for this concentration when the

investment and operational cost and the price of the raw material were varied showed that the probability that the production cost

could be lower than 0.90 or 1.0 US$/kg was 61% or 91%, respectively. Ó 2000 Elsevier Science Ltd. All rights reserved.

Keywords: Lactic acid; Economy; Process; Wheat ¯our

1. Introduction

The production of lactic acid has attracted a greatdeal of interest due to its potential use as a raw materialin the production of the biodegradable polymerpoly(lactic acid). The world production of lactic acid isapproximately 50,000 tonnes/year (Datta and Tsai,1997) and the commercial price of lactic acid rangesfrom 1.40 US$/kg for 50% to 1.90 US$/kg for 88% foodgrade lactic acid (Chem. Mark. Rep., 1999). The fer-mentative production of lactic acid is interesting due tothe prospect of using cheap polysaccharide raw materi-als such as starch or cellulose. However, the productionof lactic acid, using this process, su�ers from a numberof drawbacks. The raw material must be exposed tosome form of pre-treatment before fermentation toproduce a suitable fermentation medium. Additionally,lactic acid fermentation is product inhibited (Goncßalveset al., 1991; �Akerberg et al., 1998) and the production oflactic acid at high concentrations is thus not favourable.

Since lactic acid is acidifying, the fermentation step re-quires pH control and the addition of a titrating agent,e.g. sodium hydroxide, ammonium hydroxide or calci-um carbonate, resulting in lactate formation. In the re-conversion of lactate to lactic acid, undesired by-prod-ucts may be produced, e.g. calcium sulphate (Dattaet al., 1995). At high temperatures and high concentra-tions, lactic acid is polymerised (Mellan, 1977; Cockremand Johnson, 1993), which complicates the recovery ofthe lactic acid by conventional techniques such asevaporation. In order to develop an economicallycompetitive process with a low degree of by-productformation, the optimal operating conditions must beidenti®ed for all process steps. An economical evalua-tion of the process can be used to identify the bottle-necks and attention can then be focused on costly partsof the process, which must be improved.

Earlier economical studies of the fermentative pro-duction of lactic acid have resulted in a range of pro-duction costs of the ®nal product. Most of these studieshave only considered parts of the process, for example,the upstream processing (Vick Roy, 1983; Timmer andKromkamp, 1994; Kaufman et al., 1995) or the down-stream processing (Kramer and Al-Samadi, 1995;

Bioresource Technology 75 (2000) 119±126

* Corresponding author. Tel.: +46-46-222-82-97; fax: +46-46-222-45-

26.

E-mail address: [email protected] (G. Zacchi).

0960-8524/00/$ - see front matter Ó 2000 Elsevier Science Ltd. All rights reserved.

PII: S 0 9 6 0 - 8 5 2 4 ( 0 0 ) 0 0 0 5 7 - 2

Tejayadi and Cheryan, 1995), and in some investigations,the ®nal product consisted mainly of lactate rather thanlactic acid (Vick Roy, 1983; Millis, 1993; Timmer andKromkamp, 1994). In these studies, the lactic acid pro-duction cost ranged from 0.10 to 2.00 US$/kg dependingon how much of the process was considered and thequality of the ®nal product. A few complete processeshave been suggested where lactic acid, at concentrationsof 50% (Cable and Sitnai, 1971) and 82% (Datta et al.,1995), could be produced at a cost of 0.26 AU$/kg and0.55 US$/kg, respectively. For comparison, it can bementioned that the synthetic production of lactic acidcosts between 1.30 and 1.40 US$/kg (Millis, 1993).

In the present study, the production cost of 70% lacticacid from whole-wheat ¯our was investigated for aprocess assumed to consist of the following processsteps. The starch in the ¯our is enzymatically hydrolysedto form glucose in two steps: liquefaction and saccha-ri®cation. The glucose is converted to lactic acid in thefermentation step, using components from the wheat¯our as nutrients. The ¯our and bacteria are removed bycentrifugation and the ¯our proteins as well as the en-zymes are removed by ultra®ltration after fermentation.The fermentation step is performed at a controlled pH,producing mainly lactate. Using water-splitting electro-dialysis, the lactate is converted to lactic acid, and so-dium hydroxide is produced as a by-product. The lacticacid is concentrated by evaporation below atmosphericpressure.

The lactic acid production cost was evaluated usingkinetic models for the sacchari®cation (�Akerberg et al.,2000) and fermentation (�Akerberg et al., 1998) steps

developed earlier and using data from other experi-mental studies performed previously (�Akerberg andZacchi, 1999). The bottlenecks were identi®ed and dif-ferent process alternatives were evaluated in order to®nd the optimal operating conditions for the wholeprocess. Examples of such process alternatives are: therecovery of wheat ¯our as fodder, integrating the sac-chari®cation and fermentation steps, recovery of sodiumhydroxide from the electrodialysis step, and fermenta-tion at a lower pH.

2. Methods

2.1. The process (Process 1)

In the basic process, (Fig. 1, Process 1), no processsteps are integrated and the two hydrolysis steps, theliquefaction and the sacchari®cation, as well as the fer-mentation step are performed as batchwise processes.The di�erent process steps and assumptions made arepresented below.

2.1.1. Raw materialWheat consisting of 65% starch, is ground to ¯our

and mixed with water to a concentration correspondingto 50±180 g glucose/l after total hydrolysis of the starch.

2.1.2. LiquefactionIn this ®rst hydrolysis step, the starch in the ¯our is

enzymatically degraded to oligosaccharides at a hightemperature, 90°C, using the commercial thermostable

Fig. 1. Schematic ¯owsheet for Process 1.

120 C. AÊ kerberg, G. Zacchi / Bioresource Technology 75 (2000) 119±126

a-amylase, Termamyl 120L (Novo Nordisk, Bagsvñrd,Denmark) with an enzyme activity of 120 KNU/g,(Novo Nordisk product sheet, 1990) at a concentrationof 200 ll/kg wheat ¯our. The step was assumed to beperformed batchwise and thus, a residence tank wasrequired. The recommended reaction time for liquefac-tion is 30±60 min. (Novo Nordisk product sheet, 1990).Therefore, a residence time of 45 min. was assumed inthe present study and from this value, the workingvolume, which was assumed to be 80% of the reactortank volume, was calculated. The residence tank wasassumed to be 20% greater than the reactor tank. Thepower consumption for agitation was calculated usingpower correlations and scale-up conditions (Geankoplis,1993) assuming that the agitator was a ¯at six-bladeturbine, and that the agitation speed should be scaled-upfrom a speed of 300 rpm in a 1 l fermentor.

2.1.3. Sacchari®cationIn the second hydrolysis step, the oligosaccharides

produced in the liquefaction step are further hydrolysedto produce glucose using a commercial enzyme mixture(SAN Super 240L, Novo Nordisk) consisting of an a-amylase, amyloglucosidase and a protease. The con-centration of the enzyme was assumed to be 7 ml/kgstarch. As in the liquefaction step, it was assumed thatthis step was performed batchwise, thus requiring aresidence tank. The time required for complete conver-sion of oligosaccharides to glucose was determined bysimulation assuming batchwise operation and using akinetic model for the sacchari®cation step developedearlier (�Akerberg et al., 2000). It was assumed that thepH was 5 and the temperature 55°C in the sacchari®-cation step since these are optimal conditions for sac-chari®cation (Novo Nordisk product sheet, 1991). Fromthese calculations, the working volume, which was as-sumed to be 80% of the reactor volume, was calculated.The residence tank volume was assumed to be 20%greater than the fermentor volume. The power require-ment and scale-up of the agitator were calculated as forthe liquefaction step.

2.1.4. FermentationIn the fermentation step, the glucose produced in the

sacchari®cation step is converted to lactic acid usingLactococcus lactis ssp lactis ATCC 19435 utilisingcomponents from the wheat ¯our as nutrients. Theworking volume, which was assumed to be 80% of thefermentor volume, was determined by simulation as-suming fermentation to be performed batchwise using akinetic model developed previously (�Akerberg et al.,1998) and considering the extra time required for ®llingand emptying the fermentor. In Process 1, fermentationwas assumed to be performed at pH 6 and 30°C sincethese are optimal conditions for Lactococcus lactis(�Akerberg et al., 1998) and as a batchwise process. The

cost of a residence tank 20% larger than the fermentor,and the power requirement for the agitator as forthe liquefaction step were included in the economicevaluation.

2.1.5. Inoculum fermentationWhen batchwise fermentation was considered, an

extra process step for inoculum production was includedin the process. It was assumed that the fermentationmedium for inoculum preparation was totally hydroly-sed wheat ¯our and that the inoculum was produced togive an initial cell mass of 0.02 g/l in the fermentationstep.

2.1.6. CentrifugationThe suspended ¯our particles are removed by cen-

trifugation after fermentation. It was assumed that adecanter centrifuge was used and that the requiredpower was 40 kWh/tonne solid material discharged(Perry and Chilton, 1973).

2.1.7. Ultra®ltrationProteins in solution are removed by ultra®ltration.

The retentate stream containing proteins was assumedto be 20% of the feed stream into the ultra®ltration unitand the ¯ux was assumed to be 20 l/m2 h.

2.1.8. ElectrodialysisThe protein-free product stream is transported to a

water-splitting electrodialysis unit. The sodium lactateproduced in the fermentation step is converted here tolactic acid, and sodium hydroxide is produced by amodule containing two bipolar membranes splittingwater into hydrogen and hydroxide ions, and a cationexchange membrane transporting sodium ions from thelactate compartment to the compartment where sodiumhydroxide is produced. The transport of ions, J, over amembrane has been described by Boniardi et al. (1997)

J � iAF

g; �1�

where i is the current density (A/m2), A is the membranearea (m2), F is FaradayÕs constant (A h/mol), and g is thecurrent e�ciency. According to the membrane supplier,the current e�ciency for the membranes, which wasused in the present evaluation, is greater than 0.94 andthus, this value was used for g. A water transport of0.043 l H20/mol Na�was assumed (Lee et al., 1998). Inthe electrodialysis module, 90±95% of the lactate can beconverted to lactic acid (Bar and Byszewski, 1996). Inthe present study, a conversion of 95% was assumed.The number of stacks required was calculated assumingthat the membrane area was 0.4 m2 and the currentdensity 1080 A/m2. The power consumption was calcu-lated from the voltage drop over the membranes and thecompartments. The voltage drop was assumed to be

C. AÊ kerberg, G. Zacchi / Bioresource Technology 75 (2000) 119±126 121

4.86 V/cell over the membranes and 6.62 V/cell over thecompartments except over the compartment containingsodium lactate where the voltage drop was determinedfrom the conductivity of sodium lactate, as described byKohlrauschÕs law (Atkins, 1986). The voltage drop inthe boundary region of the membrane was neglected.

2.1.9. Concentration of lactic acid by evaporationAfter electrodialysis, the lactic acid is concentrated to

70% (w/w) in the evaporation step. Evaporation wasassumed to be performed co-currently in three stepsbelow atmospheric pressure due to the polymerisationproperty of lactic acid. The evaporator temperatures forthe three steps were assumed to be 89°C, 76°C and 56°C.It was assumed that no lactic acid entered the vapourphase after separation and that the increase in theboiling point of water due to the presence of lactic acidcould be calculated from

DT � 5:32w; �2�where DT is the increase in boiling point (K) and w is themass fraction of lactic acid (g lactic acid/g solution)(�Akerberg and Zacchi, 1999). A condenser was also in-cluded in the calculations after the last evaporator step.Steam at 1 bar was assumed to be used for heating the®rst evaporator. The heat transfer coe�cients for thethree evaporators and the condenser were assumed to be4.8, 3.0, 0.9 and 3.0 kW/m2 K, respectively.

2.1.10. Heat exchangersThe two heat exchanger operations requiring most

energy were considered in the economical evaluation;one used for heating the wheat ¯our suspension from20°C to 90°C before the liquefaction step and the otherfor heating the lactic acid solution from 20°C to 89°Cbefore the evaporator step. In both cases, steam at 1 barwas assumed to be the heating medium. The heattransfer coe�cients for the heat exchangers heating thewheat ¯our suspension and the lactic acid solution wereassumed to be 0.5 and 1.5 kW/m2 K, respectively.

2.2. Alternative process con®gurations

A number of alternative processes, Processes 2±6,were studied in order to investigate how integration andphysical properties in¯uenced the production cost oflactic acid.

2.2.1. Fodder production (Process 2)In Process 2, the proteins from the ultra®ltration step

are concentrated by evaporation in one step and mixedwith the remaining wheat ¯our from the centrifugationstep. This mixture is then dried and the ®nal productsold as fodder. Evaporation was assumed to be per-formed at 30°C, due to the presence of temperature-sensitive proteins, using steam at a temperature of 50°C

as the heating medium. The heat transfer coe�cient forthe evaporator unit was assumed to be 1.5 kW/m2 K. Itwas assumed that the mixture of ¯our and proteins wasdried using a drum dryer to a ®nal dryness of 90%. Thewater-removing capacity and the area of the drum wereassumed to be 0.01 kg/m2s and 35 m2, respectively(Coulson and Richardson, 1991).

2.2.2. Simultaneous sacchari®cation and fermentation(Process 3)

In Process 3, the sacchari®cation and fermentationsteps were integrated i.e., performed simultaneously(SSF). This integration can reduce the enzyme require-ment. Due to inhibitory e�ects of the glucose as a productin the sacchari®cation step and as a substrate in the fer-mentation step, the lactic acid production rate increaseswhen the two process steps are integrated. Furthermore,the number of reaction tanks can be reduced. However, ithas been observed that the release of nutrients due toenzymatic hydrolysis is critical and may reduce theproduction rate in the fermentation step (Hofvendahlet al., 1999). In the present study, the reactor volume wasdetermined using the previously developed kinetic mod-els for the sacchari®cation and fermentation stepsassuming that the level of nutrients was not limiting, andthat fermentation was performed at pH 6 and 30°C.

2.2.3. Recycling sodium hydroxide from the electrodial-ysis step (Process 4)

In Process 4, the sodium hydroxide produced in theelectrodialysis step is recycled to the fermentor. How-ever, it must ®rst be concentrated in order not to dilutethe fermentation broth. The sodium lactate solution isthus concentrated before the electrodialysis step byevaporation at 100°C in one step to a concentrationcorresponding to 200 g/l sodium hydroxide. The heattransfer coe�cient for the evaporator unit was assumedto be 1.5 kW/m2 K.

2.2.4. Lower pH values in the fermentation step (Process5)

In Process 5, a reduced pH value is used in the fer-mentation step. Reducing the pH reduces the amount ofsodium hydroxide required. A lower pH also reduces thecost of the electrodialysis step. On the other hand, thelactic acid production rate decreases when the pH islowered from 6 to 4 (�Akerberg et al., 1998). The fer-mentation volume was determined using the fermenta-tion model developed earlier at pH 4 and pH 5(�Akerberg et al., 1998).

2.2.5. Continuous fermentation with cell recycling (Pro-cess 6)

An alternative process con®guration (Process 6), inwhich fermentation is performed continuously with cellseparation by micro®ltration and recycling of the cells to


the fermentor, was evaluated. The wheat ¯our is re-moved by centrifugation before the fermentation step.Since cell growth is totally inhibited at high lactic acidconcentrations, fermentation was assumed to be per-formed in two steps. To maintain a constant cell con-centration, a bleed stream (1% of the stream from thecentrifuge) was assumed. The membrane was assumedto be ceramic with a ¯ux of 80 l/m2 h.

2.3. Economy

The cost of major equipment was obtained fromsuppliers or estimates (Ulrich, 1984). Cost data, notconsidering costs for e.g., installation and instrumenta-tion, were multiplied by a Lang factor including extracosts for piping, instrumentation, buildings, facilities,engineering, construction, size, and unexpected costs(Peters and Timmerhaus, 1991). The value of the Langfactor was assumed to be 3.6. A pay-o� time of 15 yearsand an interest rate of 5% were assumed, giving an an-nuity present worth factor of 0.0963 (Peters and Tim-merhaus, 1991). For scale-up of equipment costs, theequation

c1 � c0

Cap1

Cap0

� �n

�3�

was used where c represents the equipment cost, Cap isthe capacity of an operation, n is the scale-up factor andthe subscripts 0 and 1 represent the reference unit andunit under study, respectively. For smaller reactors andtanks, a scale-up factor of 0.67 was used and for fer-mentors, larger tanks, membranes, heat exchangers, andevaporators, n was assumed to be 1. The cost data andcapacities in Eq. (3) for di�erent unit operations arepresented in Table 1, and operational costs are pre-sented in Table 2. An exchange rate of 8.20 SEK/US$was used for the conversion of Swedish cost data to

US$. It was assumed that the plant worked 7200 h/yearand that four technicians worked in ®ve shifts at theplant.

3. Results

The production cost of 70% lactic acid, using thecon®guration in Process 1, was evaluated. In all calcu-lations, the computer program BioProcSim (�Akerberg,2000) was used. A lactic acid production capacity of3� 104 tonnes/year was assumed. The production costand the amount of product formed per unit mass rawmaterial were evaluated for di�erent wheat ¯our con-centrations corresponding to 50, 70, 90, 130 and 180 g/lglucose after the sacchari®cation step and are presentedin Table 3. The amount of lactic acid produced per unitmass raw material decreases with increasing wheat ¯ourconcentration, while the cost decreases with increasingwheat ¯our concentration up to 70 g glucose/l. The in-vestment and operational costs for this concentration,

Table 1

Calculation of investment costs, including installation and instrumentation, using Eq. (3)

Unit Capacity unit (Cap1) c0 (103 US$) Cap0 n

Grinder Wheat ¯our ¯ow rate 49.6 20 kg/s 0.65

Reactor tank Volume 254 628 m3 0.65

Residence tank Volume 254 628 m3 0.65

Agitator Power 49.6 50 kW 0.65

Centrifuge Wheat ¯our ¯ow 124 5 kg/s 0.65

Fermentor Volume 1077 1200 m3 0.65

Micro®lter membrane and module Membrane area 26.3 1 m2 1

Ultra®lter membrane and module Membrane area (m2) 6.58 1 m2 1

Dryer Number of drums 1230 1 1

Electrodialysis unit Number of cells 7.02 1 1

Electrode Number of cells 3.51 1 1

Heat exchangers Heat exchanger area 0.95±1.45a 1 m2 1

Separatorsb 54.9 1 1

a The di�erent values refer to heat exchangers of di�erent materials, using di�erent gasket materials.b The separators are used to separate liquid and vapour in the evaporators.

Table 2

Operational costs

Item Operational cost

Labour 48,800 US$/employeeáyear

Process water/cooling water 0.0122 US$/m3

Power 0.0366 US$/kWh

Steam 0.0220 US$/kg

Wheat ¯our 0.134 US$/kga

0.0915 US$/kgb

Enzyme (Termamyl) 3.54 US$/l

Enzyme (SAN Super) 8.05 US$/l

Sodium hydroxide (40%) 0.292 US$/kg

Cleaning of membrane 0.122 US$/m3 permeate

Ultra®ltration membrane 122 US$/m2 year

Electrodialysis membrane 439 US$/cell year

a No fodder production.b Fodder production, the price is calculated as the di�erence between

the wheat cost and the price obtained for the fodder.


distributed over the di�erent parts of the process, arepresented in Fig. 2.

It was observed that the operational cost contributedabout 80% to the total cost. The major operational costsare the cost of raw material, sacchari®cation, fermen-tation, and electrodialysis. To reduce these costs, theprocess can be integrated or the pH value lowered ac-cording to Processes 2±6. The lactic acid production costfor these processes was evaluated and is presented inFig. 3.

To reduce the cost of the raw material, the recoveryof wheat ¯our and proteins for the production of fodderwas evaluated (Fig. 3, Process 2,). It was found to beadvantageous to produce fodder only when high wheat¯our concentrations are used.

The major contributor to the operational cost in thesacchari®cation step is the cost of the enzyme mixture.This can be reduced by integrating the sacchari®cationand fermentation steps (Process 3). The cost was eval-uated assuming that the level of nutrients was not lim-iting (Fig. 3, Process 3). It was observed that the costcould be reduced by 0.02±0.12 US$/kg using SSF.

The cost of sodium hydroxide dominates the opera-tional cost in the fermentation step and recycling of

sodium hydroxide from the electrodialysis step to thefermentor was thus evaluated (Fig. 3, Process 4). It wasfound that the cost could be reduced by as much as 0.26US$/kg for the highest wheat ¯our concentration. An-other way of reducing the sodium hydroxide consump-tion is to perform fermentation at a lower pH. Theproduction cost was evaluated at pH 4, 5 and 6 andFig. 3 shows that the cost decreases with decreasing pH.It has been reported (Vick Roy, 1983; Timmer andKromkamp, 1994) that the continuous production oflactic acid with cell recycling is economically favourable.The lactic acid production cost was thus estimated forthis process con®guration and compared with that inProcess 1 (Fig. 3, Process 6). However, this process al-ternative did not reduce the cost, based on the data usedin this study.

Numerical optimisation with respect to the wheat¯our concentration was performed for Process 4, whichwas considered the best process con®guration. Thisshowed that the optimal wheat ¯our concentrationcorresponded to a glucose concentration of 116 g/l, re-sulting in a lactic acid production cost of 0.88 US$/kg.

To investigate the sensitivity of the cost data, aMonte Carlo simulation was performed. The wheatprice and investment and operational costs were as-sumed to vary with a Gaussian distribution with stan-dard deviations of 10, 20 and 20%, respectively. Thesimulation showed that the probabilities that the pro-duction cost would be lower than 0.90 and 1.0 US$/kgare 61% and 91%, respectively (Fig. 4).

4. Discussion

The production cost of lactic acid with a concentra-tion of 70% (w/w) was evaluated for di�erent process

Table 3

The lactic acid production cost at di�erent wheat ¯our concentrations

for Process 1

Glucose concen-

tration (g/l)

Cost (US$/ kg

lactic acida)

Amount of 70%

lactic acid pro-

duced per unit

mass raw material

(kg/kg)

50 1.10 0.64

70 1.06 0.58

90 1.08 0.54

130 1.10 0.47

180 1.15 0.39

a Lactic acid with a concentration of 70%.

Fig. 2. The investment (h) and operational costs (j) for Process 1

using a wheat ¯our concentration corresponding to 70 g/l. Hydrolysis

comprises the liquefaction and sacchari®cation steps, separation, the

centrifugation and ultra®ltration steps and recovery comprises the

electrodialysis and evaporator steps.

Fig. 3. The lactic acid production cost (US$/kg) for di�erent glucose

concentrations (g/l) for di�erent process con®gurations; Process 1 (r),

fodder production (Process 2) (m), simultaneous sacchari®cation and

fermentation (Process 3) (*), recovery and recycling of sodium hy-

droxide (Process 4) (´), pH 4 (Process 5) (d), pH 5 (Process 5) (+) and

continuous fermentation (Process 6) (j).


con®gurations. The optimum wheat ¯our concentrationwas determined and a Monte Carlo simulation wasperformed to investigate the sensitivity of cost data. Itshould be observed that only the actual process costsand not the cost of, for example, transportation andstorage were considered in the cost evaluation. Costevaluation is, however, a valuable tool in identifyingbottlenecks in the process and for the comparison ofdi�erent processes.

Fig. 3 shows that it is only advantageous to recoverwheat ¯our and proteins as fodder (Process 2) at thehighest wheat ¯our concentration studied since the re-covery cost, mainly due to drying, was higher than theexpected income from the fodder at the lower con-centrations. It is also uncertain whether fodder con-taining large amounts of lactic acid is an attractiveproduct.To reduce the raw material cost, an alternativeraw material must be found, either containing both acarbon source and nutrients, as does wheat, or a car-bon source which can be complemented with a cheapnutrient.

Integration of the sacchari®cation and fermentationsteps reduces the lactic acid production cost slightly(Fig. 3). The cost of the enzyme mixture is decreased andthe extra tanks for the sacchari®cation step can beeliminated. The cost evaluations for simultaneous sac-chari®cation and fermentation were performed assum-ing that the amount of nutrients in the wheat ¯our issu�cient for the fermentation to be performed at max-imum level i.e., the addition of extra nutrients will notincrease the product formation rate. Therefore, no costsfor extra nutrients were included in the economicalevaluation. It has been observed that the amount ofnutrients is su�cient when a wheat ¯our concentrationcorresponding to 180 g/l glucose and an enzyme con-centration of 7 ml/kg starch are used while at 90 g/l, therelease of nutrients due to enzymatic hydrolysis is notsu�cient and an extra nutrient source is required(Hofvendahl et al., 1999). For simultaneous sacchari®-cation and fermentation to be advantageous at a lower

enzyme concentration, or at lower wheat ¯our concen-trations, a cheap nutrient must be found.

The recovery of sodium hydroxide produced in theelectrodialysis step and recycling of this to the fermentorreduces the lactic acid production cost considerably(Fig. 3). Since the product stream must be concentratedby evaporation before the electrodialysis step in order toproduce sodium hydroxide with a su�ciently high con-centration, the cost reduction increases with increasingwheat ¯our concentration.

The cost evaluations indicated that performing fer-mentation at a lower pH is favourable (Fig. 3). Lower-ing the pH decreases the cost of sodium hydroxide in thefermentation step, as well as the cost of conversion oflactate in the electrodialysis step. On the other hand,since the product formation rate decreases with pH(�Akerberg et al., 1998), the investment cost for the fer-mentor step increases. However, the bene®ts of reducedsodium hydroxide consumption and cheaper electrodi-alysis are greater than the increase in investment cost. Itshould be observed that the kinetic model for pH 4 andpH 5 is based on fermentation experiments performedover a short time period, producing only low concen-trations of lactic acid (�Akerberg et al., 1998). The valueof the fermentor volume is therefore uncertain.

Continuous fermentation is not favourable comparedwith batchwise fermentation (Fig. 3, Process 6). Cellgrowth is totally inhibited at lactic acid concentrationshigher than 73 g/l at pH 6 (�Akerberg et al., 1998). Evenif several fermentors in a cascade are used, at least onemust operate near this concentration where the rate ofcell growth and thus, the product formation rate, is verylow. To compensate for the long fermentation times,large fermentation volumes are required increasing theinvestment cost for fermentation when the wheat ¯ourconcentration is high. For lower wheat ¯our concen-trations where the lactic acid concentration after thefermentation step is lower than 73 g/l, the required fer-mentor volume will be smaller for continuous fermen-tation (Process 6) than for batchwise fermentation(Process 1). However, for continuous fermentation, amicro®ltration unit associated with a high investmentcost is required for cell separation. This results in ahigher lactic acid production cost for all wheat ¯ourconcentrations for Process 6.

The optimum wheat ¯our concentration for Process 4corresponds to 116 g glucose/l. The cost of separationand recovery decreases and the cost of fermentation andenzymatic hydrolysis increases with increasing wheat¯our concentration. The Monte Carlo simulationshowed that, by assuming that the wheat price, invest-ment and operational costs vary with a Gaussian dis-tribution with standard deviations of 10%, 20% and 20%respectively, the probabilities that the price would belower than 0.90 and 1.0 US$/kg are 61% and 91%, re-spectively (Fig. 4). The investment cost does not in¯u-

Fig. 4. Probability of various production costs of lactic acid, obtained

by Monte Carlo simulation.


ence the total cost to any large degree and the variationin this cost will not a�ect the total production cost asmuch as the operational cost and the wheat price.

To investigate whether a reduction in pH in the fer-mentation step is as advantageous as this study indi-cates, further experimental studies must be performed.Since the in¯uence of nutrients on simultaneous sac-chari®cation and fermentation has not been investigatedin detail and the economical evaluation is thus uncer-tain, this might be worth studying in more detail. Itwould also be of interest to investigate alternative rawmaterials and nutrients with a view to reducing the cost.

Acknowledgements

We wish to thank the Swedish National Board forIndustrial and Technical Development for their ®nancialsupport.

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An economic evaluation of the fermentative production of lactic acid