Biotechnol. Prog. 1991, 7, 185-194
185
Acetone-Butanol-Ethanol (ABE) Fermentation and Simultaneous Separation in a Trickle Bed Reactor? Chang-Ho Park,$ Martin R. Okos,*and Phillip C. Wankat Departments of Chemical and Agricultural Engineering, Purdue University, West Lafayette, Indiana 47907
A new in situ product separation technique using gas stripping was applied t o acetonebutanol-ethanol (ABE) fermentation in a gas-phase-continuous immobilized cell trickle bed reactor. Both numerical and experimental studies showed that glucose conversion could be improved when the inhibitory compounds were stripped by fermentation gas. Solvents were stripped preferentially from the fermentation broth, and butanol removal was as efficient as acetone removal in spite of butanol's high boiling point (117 "C) because of butanol's high volatility a t fermentation concentrations. Since most of the butanol was removed by gas stripping, organic acids played major roles among inhibitory products. Numerical analysis indicated that improvement of glucose conversion by gas stripping was more pronounced a t higher gas flow rates, with stronger inhibition kinetics, and under vacuum conditions. Numerical calculation predicted that glucose concentrations higher than 60g/L could be converted, but this could not be shown experimentally because of increased degeneration.
Introduction Various kinds of in situ product separation methods have been proposed to reduce product inhibition and to improve solvent productivity during ABE fermentation in liquid-phase-continuous reactors. These methods include adsorption onto active carbon (Yamasaki et al., 1958), solvent extraction using biocompatible organic solvents (Ishii et al., 1985; Taya et al., 1985; Roeffler et al., 1988; Eckert and Schugerl, 1987;Evans and Wang, 1988),reverse osmosis (Garcia et al., 1986), pervaporation (Groot et al., 1984a,b; Matsumura and Kataoka, 1987), and ultrafiltration (Ferras et al., 1986). In this study a gas-phasecontinuous immobilized cell reactor-separator concept (ICRS) developed for ethanol fermentation (Dale et al., 1985a,b) was extended to ABE fermentation for butanol production. Dale et al. (1985a,b) improved ethanol productivity in a gas-phase-continuous immobilized cell reactor-separator (ICRS). Cells were grown and entrapped in porous matrices (sponges) packed in columns and were in direct contact with the liquid and gas flow. Products in the liquid stream could be removed from the reactor by cocurrent or countercurrent stripping with circulating gases that were produced during the fermentation. The gas stream was stripped of volatile products before it was recycled. The underlying thermodynamic driving forces of this separation are a low boiling point of ethanol (78.5 "C) and its highvolatility. ABE fermentation produces solvents (acetone, butanol, and ethanol) and organic acids (acetic and butyric acids). Under normal operating conditions butanol is the major toxic compound. In spite of the high boiling point of butanol (117 OC) the thermodynamic equilibrium data for the 1-butanol-water t Contribution No. 17 468 of the Minnesota Agricultural Experiment Station. * To whom correspondence should be addressed. t Current address: Department of Agricultural Engineering, Institute for Advanced Studies in Biological Process Technology, University of Minnesota, St. Paul, MN 55108.
system reveal that the activity coefficient of 1-butanol is very high a t the dilute concentrations observed during fermentation (Stockhardt and Hull, 1931). Thus, it should be possible to strip butanol from the fermentation broth. Studies of ABE fermentation in an immobilized cell, trickle bed reactor without in situ product separation have been successfully performed (Park et al., 1989). The overall objective of this study is to predict numerically the effects of the various system variables on the improvement of glucose conversion with in situ separation and experimentally confirm the model prediction.
Experimental Procedures Clostridium acetobutylicum ATCC 824 was obtained from the American Type Culture Collection. After heat activation (75 OC, 2 min) and subsequent cooling, cells were transferred twice in RCM (reinforced clostridial medium) preculture medium before they were inoculated into the reactor. Continuous experiments were performed in two 31.5mm id., 40.6- and 61.0-cm-longglass columns (the enricher and the stripper, respectively) in a completely anaerobic environment. The columns were packed with three polyester sponge strips, which were fixed by fabricated iron wire screens. The cells were adsorbed on the surface or entrapped in the pores of the polyester sponge. The reactor columns and tubing were steam-sterilized. After assembly the apparatus was disinfected by flowing 70 5% ethyl alcohol through the system for 24 h. The columns were then washed with sterilized distilled water. Favorable cell growth conditions were ensured by supplying nutrient medium overnight before inoculation. Polyester sponge strips that were used to entrap the microorganisms were 3-4 mm thick, 30.5 and 43.2 cm long and weighed 3.58 and 7.82 g in the first and the second column, respectively. The wetted volumes of the first and second column were 48.5 and 67.3 mL, respectively. Nutrient medium and glucose solution were prepared a t double strength and mixed on-line to prevent a browning
8756-7938/91/3007-0185$02.50/0 0 1991 American Chemical Society and American Institute of Chemical Engineers
Biotechnol. Prog., 1991, Vol. 7, No. 2
186
Cocurrent operation
Glucose
I
stage j
Countercurrent operation j-1 I u I/
i , j-1
I
i
stage j
Figure 2. Scheme of equilibrium staged model.
Figure 1. Flow diagram of an immobilized cell trickle bed reactor for ABE fermentation with gas stripping; single line (-) is liquid, double line (=) is gas. reaction. The composition of the nutrient medium included the following (in grams per liter): yeast extract, 10;peptone, 10;KH2P04,0.4; Na2HP04,0.6; MgS0~7H20, 0.2; FeC1~6H20,0.01; NH4C1, 0.8; cysteine, 0.5; and 0.5 mL of trace elements. The trace element solution contained the following (in grams per liter): CaC12.2H20,0.66; ZnS04-7H20,0.18;CuS04.5H20,0.16; MnS0~4H20,0.16; and CoC1~6Hz0,0.18. Free oxygen was removed from the media by boiling the media while sparging with nitrogen. To determine the decrease in redox potential, 0.5 g/L cysteine hydrochloride (Sigma) and 0.0001% (w/ v) resazurine (Kodak) were subsequently added to boiling medium. A decrease in redox potential was noticed by the color change of resazurine from pink to colorless. Media were autoclaved at 121 OC, 15 psig for 20 min and cooled down. Glucose concentration was 15-20 g/L after on-line mixing during the cell growth period (ICR). During the ICRS experiment glucose concentrations were 20, 60, or 80 g/L after on-line mixing with nutrient. Feed glucose concentrations in the next part of the paper refer to concentrations after on-line mixing. Absorber water with 0 , l T resazurine redox indicator was boiled in an autoclave for 15 min, and cysteine hydrochloride ( 5 g/L) was added while the water temperature was above 90 OC. The water was autoclaved for another 10 min and cooled down. Nitrogen was used to keep the air (oxygen) from the headspace of the medium and absorber water bottles. The reactor operated in either the ICR or ICRS mode. During ICR experiments the liquid feed was passed through the columns without gas circulation and the reactor effluent was displaced by the generated fermentation gas. The reactor effluent line was U-shaped to prevent air (oxygen) backflow into the reactor. The ICR unit was converted to an ICRS unit by adding gas lines, a gas pump, and a product recovery unit (the absorber). The ICRS unit was flushed with N2 for 5 days before start-
up. During ICRS operation, the circulating fermentation gas strips the products from the liquid stream cocurrently in the enricher and countercurrently in the stripper and is sent to the product recovery unit (the absorber) (Figure 1). After product removal the fermentation gas is circulated back to the fermentor by a gas pump. Product effluent was collected by a peristaltic pump. The flow rate was measured as collected volume in the effluent bottle. Glucose consumption and product yield were computed considering water evaporation in the fermentation section. The absorber water effluent line was U-shaped to prevent oxygen flow through the absorber effluent collecting bottle. Total circulating gas flow rate was 15 L/min. About two-thirds of the total gas was supplied to the stripper. Reactor temperature was maintained at 36 "C. The absorber was operated at 22-23 "C.
Mathematical Model Simultaneous fermentation and separation in packed columns is modeled by using an equilibrium stage approach. The columns are viewed as composed of equilibrium stages where the gas and liquid are thermodynamically in equilibrium. The model is limited to steadystate, one-dimensional systems (axial dependence), isothermal operation, and surface reaction (no diffusion limitation in the solid phase). The material balances are derived separately for the cocurrent and countercurrent sections (the enricher and the stripper). I t is assumed that the enricher and the stripper can be modeled as being composed of equilibrium stages (Figure 2). The lengths of the enricher and the stripper are determined by the number of stages and the height of an equilibrium stage (HETP). HETP's of the solvents and acids are assumed to be the same, and the number of stages is one of the input variables. Product yields were taken from an ICR experiment (Park, 1989), assuming that product separation does not change the yield pattern. (I) The material balance of the i component on stage j in the cocurrent enriching section is lij-*
+ uij-* + rijVs= lij + uij
Biotechnol. Prog., 1991, Vol. 7, No. 2
187
or
+
Lj..lxij-l + Vj-lyij-l+ rijV, = Ljxij V j y i j where the units of the reaction rate term ( r i j ) are moles per liter of reactor per hour. By using the vapor-liquid equilibrium relationship y ij . = - Kijxij= K IJ. . 1.W ./Lj eq 1 can be solved for lid:
(7) Activity coefficients in eq 7 are estimated with Wilson's equation. Wilson's equation for a six-component mixture is
(2)
lij-l + uij-l + rijVs (3) 1.0 + K i j V j / L j Since Vj and Lj are not known initially, eq 3 must be solved iteratively. A solution is assumed to be convergent if the differences of both Lj and Vj of two successive iteractions fall within mol/h. (11) In the countercurrent stripping section, the reaction rate a t each stage is affected by the compositions on the stages above and below. Therefore, the material balance is more complex than in the cocurrent case. The steadystate material balance of component i on stage j is modified by using the relaxation method (Dale et al., 1985a). This method solves the steady-state problem by a false transient approach by adding a false time-dependent term. The false transient material balance for component i a t stage j is liJ =
. uI dx .dt2 = 1. . + uij+l - [uij + lij] + rijVs IJ-1
(4)
where
The values of molar volume (vi and vj) in eqs 9a and 9b (in milliliters per mole) are 1-butanol,91.97; acetone, 74.05; ethanol, 58.69; acetic acid, 57.54; 1-butyric acid, 92.43; and water, 18.07. Wilson's parameters (Aij and Aji) are calculated from eqs 9a and 9b by using values of X i j - Xii and Xi, - Xj, available in the literature (Gmehling et al., 1981). From eq 8 the activity coefficient for component k is 6
In yk = - l n [ ~ x j A k j ]+ 1-
dx.. UI.1J = Lj-lxiJ-l Vj+lyij+l - Vjxij- Ljxij rijVs dt where 1 and u of eq 4 are a t a new time level t + 1. The reaction rate ( r i j ) is estimated from values at the old time level t. Uj is the false accumulation parameter, the total moles of liquid on stage j . The relaxation parameter (w) is defined as w AtlU,. Using the relationships
+
+
and assuming that K , L , and V values are independent of time level and stage, we rearrange eq 4 as a system of algebraic equations:
1
'
-(xlJVa + urjjVJ (6) L The coefficients of eq 6 form a tridiagonal matrix. The Thomas algorithm with the Gaussian elimination method is used to solve for the value of liquid mole fractions at a time level t + 1. The solution is assumed to be convergent when the difference in x values between two adjacent time intervals falls within mol/h. The tridiagonal matrix has six parameters in its elements: w, V , L , V,, Kij, and r; with w as an input variable. Volume of a stage (V,) is calculated from the average HETP and the column diameter. Values of Vand L satisfying convergencecriteria are determined in the computer program. (111) Kij and r.fj were calculated as follows. The K factor of i component on stage j is estimated by assuming that vapor phase fugacity coefficients were unity:
~
i=l
j=l
or
xiAik
6
6
(10)
CxjAij j= 1
To calculate the activity coefficients of the six-component system from eq 10,15 pairs of binary Wilson's parameters (Ai, and AjJ are required. Most of the binary parameters available in the literature are for isobaric systems (Gmehling et al., 1981; Hirata et al., 1975). In this study isothermal data are needed because the system is operated at fixed temperatures. Temperature dependence of the Wilson's parameters is expressed as
b In Aij = a + T(K) A least-squares fit is used to find the values of a and b in eq 11 by using the data in Gmehling et al. (1981). Azeotropic temperatures and compositions are another source of isothermal data, and Wilson's parameter can be calculated by a Newton-Raphson method (Walas, 1985). Isobaric data are used for the binary systems where isothermal data are not available. Isothermal data are especially scarce for binary systems including 1-butyric acid. Vapor pressures of liquid components are estimated from the Antoine equation:
B T("C)+ C where Pio is in millimeters of mercury. Vapor pressure of gas (Cod is estimated from (Dale et al., 1985a). log,, Pi" = A -
P = exp
(-1979 + 17.42) T(K)
A reaction rate expression ( r ) was developed, incorporating the effects of substrate inhibition, product inhibition, product yield, and cell density along the column. Substrate inhibition was approximated by a Monod-type expression. The concentrations of butanol, butyric acid, and acetic acid are included in the product inhibition term.
Biotechnol. Prog., 1991, Vol. 7, No. 2
Table I. Experimental Conditions of Reactor Operation
I
I
run -
0.8 0.7
Ia Ib IC Id
0.6 0.5 0.4
fC
I I 0
-
I
2
.
I
*
I
-
,
.
,
.
,
4 6 8 10 12 Butanol concentration (g/L)
.
,
14
IIa IIb IIIa IIIb
.
16
Figure 3. Effect of butanol on the normalized growth rate and initial glucose uptake rate: 0,growth rate of Moreira et al. (1981); +, growth rate of Ounine et al. (1985); 0, growth rate of Leung (1982); A, growth rate of Ishii et al. (1985); W, initial glucose uptake rate of Moreira et al. (1981); *, initial glucose uptake rate of Ounine et al. (1985); model prediction for the combined effect of butanol and butyric acid on the normalized butanol production rate (-, n = 3; - - -, n = 2 in eq 14). Assuming that immobilized cells have the same inhibition kinetics as free cells, the expression for the specific butanol production rate incorporating the combined effect of butanol, butyric acid, and acetic acid is ub [g of butanol/(g of cells.h)l = .bmax&[
1.0 -
where K , is 0.1 g/L and n is 3.0 for the reference condition (see Table IV). Other values such as 2.0, 1.5, and 1.0 are used to compare with the reference condition since the literature data for butanol inhibition are scattered (Figure 3). From the literature data on free cells, the inhibition effect of butanol on the normalized cell growth rate and initial glucose uptake rate is shown in Figure 3. Inhibition by butanol and butyric acid is expressed by a single expression (the third term in eq 14) based on Leung's observation (1982). Leung showed that when butyric acid acts together with butanol, the growth rate decreases as much as when the same amount of butanol is added. It is assumed that butanol production is completely inhibited a t a total concentration of 14 g/L (butanol plus butyric acid). Data on acetic acid inhibition on free cells is obtained from Leung (1982) and expressed as an exponential term (the fourth term in eq 14). The butanol production rate at stage j (moles of butanol per liter of reactor per hour) is
The glucose consumption rate (rg)and the generation rates of other products ( r p )in stage j are
ubDo')Yp MWp yb
rp [mol of product/(L of reactorah)] = --
(17)
The maximum specific butanol production rate (Ubmax), 0.0216 g of butanol/(g of cellssh), is estimated as follows. From ICR experiments (Park, 1989),the maximum specific glucose consumption rate of 0.0933 g of glucose/(g of
no. of mode ports 15-17 g/L ICR 1 17 g/L ICR(S)" 1 17 g/L ICRS 1 60 g/L ICRS 1
20 g/L ICR 80 g/L ICRS
1 1
IIIC IIId
20 g/L ICR 20 g/L ICRS 80 g/L ICRS 80 g/L ICR
2 2 1 1
IVa IVb IVC
20g/LICR 20 a / L ICRS 80 ijL ICRS
2 2 1
flow rate feed
abs water, mL/h
duration, days 4 2 2 7
80 or 360 80 or 330
73 67
245 240
7 7
75 or 178 55
270
5
125
2
127
200
5
202
3
128 120
4 3 8
120 50
122 259 260
ICR(S) means ICR with gas circulation.
ce1ls.h) was calculated on the basis of the total cell mass (grams of cells) and the glucose consumption rate (grams per hour). A butanol yield of 13.9% and concentrations of inhibitory products (butanol 4.68 g/L, butyric acid 3.07 g/L, and acetic acid 2.33 g/L) were obtained from an ICR experiment. When these data are applied to eq 14, the maximum specific butanol production rate (obmar) is calculated as 0.0216 g of butanol/(g of ce1ls.h). It is assumed that immobilized cells in the ICRS have the same maximum specific butanol production rate and viability as in the ICR. If the cell viability in the ICRS is higher than in the ICR, the maximum specific butanol production rate will be higher than 0.0216 g of butanol/(g of ce1ls.h). Cell density a t stage j , DG), was obtained by measuring the average cell density of 60.7 g of cells/L of reactor in the enricher and 29.9 g of cells/L of reactor in the stripper after reactor shut down. From previous ICR experiment it is known that cell density is almost the same along the enricher except at the entrance. The cell density in the stripper is almost the same along the column (Park, 1989). The volume of each stage was obtained using HETP value for the 1-butanol-water system for structured packing (21.6 cm in the stripper and 43.5 cm in the enricher). HETP is obtained in a separate experiment; the details have been described previously (Park, 1989). The yield pattern is obtained from ICR experimental results (Park, 1989): butanol = 0.139, butyric acid = 0.091, acetic acid = 0.069, acetone = 0.062, ethanol = 0.035, COz = 0.540, and hydrogen = 0.015. The yield pattern is assumed to be the same along the columns because pH dropped from above 6.0 to below 5.0 within the first 2.54 cm of the reactor column and became almost constant at pH 4.5 in the remainder of the column.
Experimental Results and Discussion The effect of gas stripping was studied experimentally. The performances of the ICRS and the ICR were compared a t low and high glucose concentrations. The reactor was operated for 60 days with four different runs (runs I-IV). For each run, cell growth was accomplished in ICR operation a t a low glucose concentration and then switched to ICRS operation. The operating conditions are shown in Table I. The two-port feed pattern (see Figure 1) improved cell growth in the second (stripper) column. There was a reactor upset after run IIb, and most of the cells in the stripper were washed out by flooding, but the cells were recharged at the beginning of run IIIa. Transient profiles of glucose consumption rate and effluent pH are shown in Figure 4. Transient profiles of production rate of each product are shown in Figure 5 . Glucose con-
Biotechnol. Prog., 1991, Vol. 7, No. 2
189
-
h
f
~~
Run11
Run1
a
RvnlU
7
Table 11. Average Experimental Glucose Consumption, Yield, and Selectivity
RunN
mZ
-6
-I.
consumed glucose, solvent acid butanol run experiment g/L yield yield selectivity PH Ia 15-17 g/L ICR 12.23 0.110 0.343 0.512 4.52 12.56 0.139 0.292 Ib 17 g/L ICR(S)a 0.559 IC 17 g/L ICRS 16.15 0.116 0.320 0.525 4.66 Id 60g/L ICRS 43.36 0.254 0.146 0.534
0
10
30 days
20
40
50
60
F i g u r e 4. Experimental profiles of p H and glucose consumption rate (g/hr) vs. time. Symbols: 0,glucose consumption rate; -, PH. Run II
-
Run 111
Run IV
X
5 -
.a
20g/L ICR 80 g/L ICRS
16.45 27.55
0.130 0.341 0.160 0.272
0.472 0.520
IIIa IIIb IIIc IIId
20 g/L 20 g/L 80 g/L 80 g/L
ICR ICRS ICRS ICR
16.35 16.43 30.36 19.63
0.118 0.110 0.203 0.170
0.318 0.229 0.218 0.240
0.491 0.566 0.558 0.502
4.21 4.25 4.23 4.26
IVa 20g/LICR IVb 20g/L ICRS IVc 80g/LICRS
15.03 12.06 23.72
0.097 0.343 0.102 0.343 0.114 0.323
0.420 0.450 0.486
4.30 4.37
a
CI
b
IIa IIb
900 800 700 600 500
ICR(S) means ICR with gas circulation.
product recovery in the absorber during ICRS operation was calculated as follows:
400
300
(y
200
f
100
a
d
;
lb
1'5
210
2;
io
315
4b
415
sb
5'5
6b
days
F i g u r e 5. Experimental profiles of production rate (grams per hour) vs time. Symbols: 0 , butanol; 0,acetone; - - -,ethanol; 0,butyric acid; A, acetic acid.
sumption rate and solvent production rate were highest during run Id and run IIIc. The effluent pH was maintained above 4.2 in all runs. Glucose Consumption, Yields, and Selectivity. Average values of consumed glucose, solvent and acid yields, and selectivity of butanol for different experiments are shown in Table 11. The highest glucose consumption was 43.36 g/L (run Id). We could not obtain glucose conversion higher than 60 g/L with product separation. Solvent and acid yields were obtained by dividing the total rate of product formation (grams per hour) by glucose consumption rate (grams per hour). The total rate of product formation is the summation of product concentration in the reactor effluent and in the absorber effluent. Solvent yield increased at higher glucose concentrations. Solvent yield of the ICRS a t 60 g/L glucose in the feed (run Id) was0.254, which is comparable to the values (0.195-0.273) obtained from a previous study of the ICR (Park, 1989). However, solvent yield from the ICRS a t 80 g/L glucose (runs IIb, IIIc and IVc) decreased and was in the range from 0.114 to 0.203, and the acid yield increased in the range from 0.218 to 0.323. Butanol selectivity was also lower than was observed during the previous ICR study (0.535-0.592) especially in runs IIb and IVc. Butanol selectivity increased after flooding (run IIIb) but dropped again during run IV. These results suggest that solvent yield and butanol selectivity decreased faster during ICRS experiments than ICR. Possible reasons for a faster decrease in solvent yield would be a faster loss of cell viability and enhanced degeneration under long-term gas circulation. Decreased butanol selectivity may result from a lower hydrogen content in the liquid phase because of gas flow. In a previous study degeneration did not decrease butanol selectivity (Park et al., 1989). Efficiency of Product Recovery. The efficiency of
{[(absorberwater flow rate) X (concn in the absorber water effluent)] + [(stripper liquid flow rate) X (concn in the stripper liquid effluent) (absorber water flow rate) X (concn in the absorber water effluent)]) X 100 (18) The results showed that product recovery could be improved when the ratio of the absorber water flow rate to the feed flow rate increased from 1.58 to 5.2. The ratio is proportional to L / G, the absorber water flow rate divided by the absorber gas flow rate. The recovery of butanol and acetone ranged from 64.5 to 87.4% and from 59.0 to 71.2%, respectively. Gas stripping separated butanol as efficiently as acetone in spite of butanol's higher boiling point (117 "C) because of butanol's higher activity a t a low concentration. Stripping of organic acids was far less efficient than stripping of solvents. The recovery of butyric and acetic acid ranged from 11.9 to 37.3% and from 6.0 to 18.3% ,respectively. Relative volatilities of butanol, acetone, ethanol, acetic acid, and butyric acid to water were 44.3, 59.5, 13.3, 1.91, and 1.51, respectively. The model predicted that product recovery percentages of butanol, acetic acid and butyric acid were 97.6%, 53.8%, and 43.7 % , respectively, assuming complete product recovery in the absorber. Gas stripping selectively removes solvents instead of organic acids. This is one of the advantages of stripping with fermentation gas since a certain level of butyric acid and a low pH are required for the initiation of solvent production (Monot et al., 1984). If any separating agent removed butyric acid preferentially, butanol production would be hampered. Comparison of Experimental ICRS and ICR Results. Simultaneous product separation during fermentation (ICRS) improved reactor performance significantly at high glucose concentrations (60 and 80 g/L), but the improvement was negligible at low glucose concentrations (17-20 g/L). A t low glucose concentrations product separation did not improve reactor performance since the concentration of an easily removable toxic product, butanol, was low before gas stripping. At 80 g/L glucose concentration, the average glucose conversion was improved by 54.7% from 19.63 to 30.36 g/L with product separation. Solvent and acid production increased by 93 %
+
1BO
and 36 71, respectively. During ICRS operation at 60 g/L glucose and 67 mL/h (run Id), the average and the highest glucose conversion values were 43.36 and 56.32 g/L, respectively. Compared to a previous ICR study (Park, 1989), the rate of glucose consumption per unit cell mass during run Id increased by 27.4% in a similar reactor. Total cell masses of this ICRS and previous ICR experiments were 24.44 and 29.47 g, respectively. Glucose conversion of the previous ICR study was 29.37 g/L at 93.6 mL/h flow rate. Performance of ICRS at High Glucose Concentrations. The effect of high glucose concentrations in the feed (60 and 80 g/L) on the ICRS performance was studied. Higher glucose consumption rate and butanol production rate could be obtained at 80 g/L as compared to 60 g/L when the new cells were charged in the stripper after flooding. However, glucose conversion did not exceed that obtained with a 60 g/L feed. The average glucose conversion decreased from 43.4 to 27.6 g/L when feed glucose concentration increased from 60 g/L a t 67 mL/h (run Id) to 80 g/L a t 55 mL/h (run IIb). The rate of glucose consumption (1.52 g/h) at 80 g/L was lower compared to 2.91 g/h a t 60 g/L (run Id). We repeated the experiment at 80g/L (runs IIIc and IVc) after cell washout in the stripper by flooding (after run IIb) and subsequent cell growth at 20 g/L glucose (runs IIIa and IVa). During run IIIc the average glucose conversion increased from 27.6 to 30.4 g/L even though the flow rate was increased from 67 to 128 mL/h. The rate of glucose consumption (3.89 g/h) was better than that obtained with the 60 g/L feed in the ICRS (run Id) (2.91 g/h). Butanol production rate was higher with a feed containing 80 g/L (run IIIc) glucose (0.44 g/h) compared to a butanol production rate of 0.39 g/h with a 60 g/L glucose feed (run Id). This increase in glucose consumption rate may be because the dead cells in the stripper were replenished with young viable cells after cell washout and subsequent cell growth. However, we could not show that glucose conversion could exceed that obtained with a 60 g/L glucose feed at a reduced flow rate. During run IVc the average glucose conversion dropped to 23.7 g/L a t a lower flow rate (50 mL/h) and feed concentration of 80 g/L glucose. The rate of glucose consumption decreased to 1.19 g/h. The causes of this drop may be due to a decreased cell viability and enhanced degeneration during ICRS operation. Reactor performance a t 80 g/L glucose based on the glucose consumption rate (grams of glucose per hour) was dependent upon the cell age and medium flow rate. However, the average reactor performance during run IIb and run IIIc was 2.70 g/h, which was 92.8% of the reactor performance with a 60 g/L glucose in the ICRS (run Id). These results suggest that substrate inhibition to immobilized C. acetobutylicum a t a feed concentration of 80 g/L is almost the same as a t 60 g/L, since product inhibition for 60 g/L (run Id) and a feed concentration of 80 g/L (runs IIb, IIIc and IVc) were almost the same (Table 111). The specific butanol production rate in the ICRS with a glucose feed concentration of 60 g/L was 0.016 g butanol/(g of cellwh), and butanol yield was 13.6%. Cell Loading Measurement. The sponge strips and the fabricated wire screen in the first and the second column were cut into five and six pieces of equal length (6.1 and 7.2 cm each), respectively. The total cell masses in the enricher and the stripper were 14.38 and 10.06 g, respectively. The cell densities of the enricher from the top to the bottom (in grams of dry cells per gram of sponge) were 3.97, 3.60, 3.25, 2.89, and 3.53 (average 3.47), and those in the stripper were 1.04, 1.45, 1.49, 0.95, 1.04, and
Biotechnol. Prog., 1991, Vol. 7, No. 2
Table 111. Experimental Concentrations of Inhibitory Products in the Effluent rate of butanol production, g/hx 55.1
run Ia Ib IC Id
butanol, acids, total, experiment g/L g/L g/L 15-17 g/L ICR 0.69 4.19 4.88 17 g/L ICR(S)a 3.86 4.71 0.85 17 g/L ICRS 4.61 4.89 0.28 60 g/L ICRS 1.83 4.82 6.65
IIa IIb
20 g/L ICR 80 g/L ICRS
1.04 0.56
5.52 6.17
6.56 6.73
IIIa IIIb IIIc IIId
20 g/L 20 g/L 80 g/L 80 g/L
ICR ICRS ICRS ICR
0.95 0.80 1.25 1.66
5.19 3.17 5.59 4.66
6.14 3.97 6.84 6.32
118.4 129.9 440.2 201.0
IVa 20g/L ICR IVb 20 g/L ICRS IVc 80 g/L ICRS
0.61 0.22 0.19
5.03 3.34 5.13
5.64 3.56 5.32
74.7 66.4
a
71.8
ICR(S) means ICR with gas circulation.
1.40 (average 1.21). The fraction of cell mass attached on the inner surface of the glass column was 2.55 g, which is 10.4% of the total cell mass.
Modeling Results and Discussion The equilibrium stage model was solved numerically. The effect of in situ product stripping on glucose conversion and product separation was studied by varying the proportion of gas/water mole ratio in the enricher and in the stripper, total gas flow rate, product inhibition kinetics, operating temperature and pressure, and inlet glucose concentration. The reference conditions are shown in Table IV. The reference conditions mean the base conditions from which the value of one variable has been changed for numerical study. Since butanol, acetone, and ethanol are easily removed and their concentrations remained below inhibitory levels, acid inhibition becomes important in an immobilized cell trickle bed reactor with product separation. It was assumed that products were completely recovered from the gas before the gas was recycled back to the fermentor. Effect of Simultaneous Product Separation. Reactor performance with and without in situ separation was compared at reference conditions in both enricher and stripper. ICR performance was determined a t a low mole ratio of the gas/liquid flow rate in both enricher and stripper (0.01/0.01 in enricher/stripper) corresponding to the amount of gas flow generated by fermentation (moles per hour). Product removal by gas stripping increased overall glucose conversion by 34.9 % : 14.05% in the enricher and 68.5% in the stripper. The profiles of glucose and inhibitory products for the ICR and the ICRS are shown in Figure 6. Glucose conversion improved better in the stripper because the difference of the inhibitory product concentrations between ICR and ICRS is greater in the stripper. The concentration profile of butanol was different from that of organic acids. Butanol concentration in the liquid phase decreased markedly in the stripper. In the case of acids the decrease was small. Theoretically and experimentally,glucose conversion higher than 60g/L cannot be accomplished without in situ separation (Park et al., 1989). With in situ product separation, the theoretical calculations predict that inlet glucose concentrations higher than 60 g/L could be processed with a longer reactor. Effect of Total Gas Flow Rate. The total gas/liquid flow rate was varied from 0.09 to 9.0 while the ratio of gas into the enricher to the stripper ratio was kept a t 2/7. As
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2
;
; 0
0 0
0
1
2
3
4
5
6
7
' i' 6
'l30
Total Gas Flow Rate (gar/water mole ratio)
Figure 7. Theoretical prediction of the effect of total gas flow rate on the consumed glucose and product concentrations in the liquid effluent of ICRS stripper (grams per liter): 0 , consumed glucose; . , butanol; 0,acetic acid; 0,butyric acid.
-
0
20 40 60 80 100 enricher stripper length of the packing (em)
120
140
Figure 6. Theoretical profiles of residual glucose and inhibitory products (numerical results): 0,without in situ separation for residual glucose, butanol, and butyric acid; 0 , with in situ separation for residual glucose, butanol, and butyric acid; 0, without in situ separationfor acetic acid;., with in situ separation for acetic acid.
0
5
10
15
20
25
Gas Proportion (strlpper/enrlcher rrllo)
Table IV. Reference Conditions for Numerical Simulation. temperature pressure feed rate feed glucose concentration total gas/water mole ratio enricher gas/water mole ratio stripper gasjwater mole ratio number of stages in the enricher number of stages in the stripper n in eq 14
36 O C 760 mmHg 67 mL/h 80 g/L 9 2 7 1 4 3
shown in Figure 7, a higher total gas/liquid flow rate increased glucose conversion and product separation. The increase in product separation was higher for butanol (more volatile component) than for acids (less volatile components). Effect of Gas Proportion. The effect of various proportions of gas flow in the enricher and in the stripper on glucose consumption and product separation was studied. Total gaslliquid flow rate was fixed a t 9 mol of gas/mol of water. As shown in Figure 8, the higher gas flow in the stripper (higher stripper/enricher ratio) reduced product loss in the stripper effluent. When 99.99 71 of the total gas was supplied to the enricher, the effluent concentrations of butanol, acetic acid, and butyric acid were 4.45,3.35, and 4.62 g/L, respectively. Butanol, acetic acid, and butyric acid lost in the effluent decreased to 0.17, 1.66, and 2.81 g/L, respectively, when 99.99%)of the gas was supplied to the stripper. Product separation improved drastically when the enricher/ stripper ratio increased from 0.001 to 0.29 and leveled out for further increases from 0.29 to 899. These results suggest the use of higher gas flows in the stripper for separation purposes even though glucose conversion decreases slightly. Use of a fraction of gas flow higher than 7 out of 9 in the stripper improves acid separation while butanol separation remains the same. Effect of Product Inhibition Kinetics. Improvement
Figure 8. Theoretical prediction of the effect of gas proportion on the consumed glucose and product concentrations in the liquid effluent of ICRS stripper (gramsper liter): 0 ,consumed glucose; ., butanol; 0,acetic acid; 0,butyric acid. in glucose conversion by in situ product separation is higher for the cells more sensitive to product inhibition. In the case of butanol, a convex upward inhibition pattern was observed (Leung, 1982; Ishii et al., 1985). This inhibition pattern suggests that butanol inhibition is not severe a t low concentrations, but it increases drastically beyond that. For a convex upward inhibition curve, the improvement by in situ separation will be small a t low butanol concentrations, but it will be large a t high butanol concentrations. Since the literature data for butanol plus butyric acid inhibition are scattered (Figure 3), four different inhibition patterns ( n = 1.0, 1.5, 2.0, and 3.0 in eq 14) were studied. A smaller value of n corresponds to stronger inhibition by butanol plus butyric acid. As shown in Figure 9, for smaller values of n, the improvement in glucose conversion by in situ separation was higher. Effect of Operating Temperature. Product separation is improved a t higher temperatures. As shown in Figure 10, glucose conversion increased from 58.72 to 68.43 g/L when the temperature increased from 25 to 42 "C because the concentration of inhibitory products in the liquid phase decreased. The production rate was assumed to be independent of temperature between 25 and 40 OC (McNeil and Kristiansen, 1985). Effect of Pressure. Operation under vacuum conditions improved glucose conversion and product separation. As shown in Figure 11, glucose consumption increased as the pressure decreased, and the improvement was higher in the enricher than in the stripper. Effect of Glucose Concentration in the Feed. As shown in Figure 12, glucose conversion decreased slightly from 64.95 to 60.96 g/L when glucose concentrations in
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I
I
I V
,.70
I
0
60-
5034.9%
.E
40-
3 0 ! 0
.
I
'
,
'
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3
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1
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Figure 9. Theoretical prediction of the effect of different inhibition (n value in eq 14) on the glucose conversion with in situ separation: 0 , ICRS (gas flow rate 2/7); 0,ICR (gas flow rate 0.01/0.01).
25
30
40
35
Temperature
45
(OC)
Figure 10. Theoretical prediction of the effect of temperature on the consumed glucose and product concentrationsin the liquid effluent of ICRS stripper (gramsper liter): 0 ,consumed glucose; m, butanol; 0,acetic acid; 0,butyric acid.
0
200
:: !e
n (power index in the inhibition kinetics)
20
f
:E
400
600
800
Pressure (mmHg)
Figure 11. Theoretical prediction of the effect of pressure on the consumed glucose and product concentrations in the liquid effluent of ICRS stripper at reference gas flow rate (grams per liter): 0 ,consumed glucose; H, butanol; 0,acetic acid; 0,butyric acid. the feed increased from 80 to 200 g/L, probably because the concentration of inhibitory products increased slightly.
Comparison of Model Predictions to Experimental Results The mathematical model and the separation characteristics of the 31.5-mm i.d. bench-scale reactor have been used to obtain predictions for run Id. This run was chosen because solvent yield, 0.254, was comparable to the value (0.273) obtained from a previous study of ICR (Park et al., 1989). The input variables are as follows: temperature, 36 "C;pressure, 760 mmHg; glucose concentration in the feed, 60 g/L; feed rate, 67 mL/h; gas flow rate, 3.3 L/min in the enricher and 11.7 L/min in the stripper. The
E Z 0 s '1
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o40
80
120
160
200
240
Feed Glucose Concentration ( g l L )
Figure 12. Theoretical prediction of the effect of feed glucose concentration on the consumed glucose and product concentrations in the liquid effluent of ICRS stripper (grams per liter): 0 , consumed glucose; H, butanol; 0,acetic acid; 0,butyric acid. number of stages in the enricher is 0.7 and in the stripper 2.0 from a McCabe-Thiele diagram of the operating line and equilibrium line (Park, 1989). For computation the number of stages in the enricher was set to 1. After the effect of one equilibrium stage was calculated, 70% of the increased vapor composition of solvents and acids was used. The remaining 30% was added back to the liquid flow. The incomplete product removal in the product recovery unit (the absorber) was also accounted for as follows: the effluent vapor composition was calculated by starting from pure gas circulation. After one calculation the vapor composition in the effluent was obtained. By use of this composition and recovery percentage in the absorber (40.8, 25.3, 28.5, 54.5, and 40.0% for butanol, acetone, ethanol, acetic acid, and butyric acid, respectively) obtained from a separate experiment (Park, 1989), the product content in the recycled gas was determined. After nine iterations the product concentrations in the stripper effluent converged within g/L. The model predicted that glucose conversion could be improved to 43.10 g/L with in situ product separation as compared to 37.56 g/L without separation. This prediction was close to the experimentally observed glucose conversion of 43.46 g/L. The profiles of the residual glucose concentration and the inhibitory product concentrations obtained by the model prediction and experiments are shown in Figure 13. Butanol concentration reached a peak in the enricher effluent and kept decreasing in the stripper. The concentration of acids kept increasing in the first part of the stripper and then decreased. In the enricher the experimental glucose consumption was more than the predicted value, probably because the viability of the enricher cells was higher than that of the stripper cells. In the model the viability of the enricher and the stripper was assumed to be the same. The maximum specific glucose consumption rate used in the model [0.0933 g of glucose/ (g of cells-h)] was based on the same cell viability in the enricher and in the stripper. The experimental value of the butanol (solvent) concentrations in the liquid effluent is higher than the predicted value.
Conclusions A new technique using gas stripping as an in situ product removal method was applied to acetone-butanol-ethanol (ABE) fermentation in an immobilized cell trickle bed reactor. The cells, liquid, and gas were in direct contact without any intervening media. Numerical studies predicted a 68.5 7% improvement in glucose conversion in the stripper by applying in situ product removal at the reference
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30 20 10 60 50 40
0
I
I
0
0
10
20
30
enricher stripper length of the packing (Inch)
Figure 13. Profiles of residual glucose and product concentrations: -, model prediction for residual glucose, butanol, and acetic acid; 0,experimentalresults for residual glucose,butanol, and acetic acid; +, model prediction for butyric acid; 0 , experimental results for butyric acid. conditions. Higher total gas flow rate and higher temperature increased glucose conversion and product separations. Higher proportions of gas flow rate in the stripper with a fixed total gas flow rate of 9 mol of gas/mol of water improved product separation but decreased glucose conversion. The improvement by in situ separation was higher for stronger inhibition kinetics. Operation under a vacuum improved glucose conversion. With in situ product removal, the model predicted that glucose concentrations higher than 60 g/L could be consumed by using longer reactors. The fermentation gas stripped solvents preferentially, and butanol removal was as efficient as acetone removal. Stripping of organic acids was far less efficient than of solvents. Experiments showed that up to 87.4% of butanol and up to 37.3 % and 18.3 % of the butyric and acetic acid were recovered by using a water absorber at a flow rate of 270 mL/h. With this removal of toxic products from the fermentor, glucose conversion improved by 33.6 and 54.7 % a t feed glucose concentrations of 60 and 80 g/L, respectively. Glucose consumption rate at a feed glucose concentration of 80 g/L increased to 3.89 g/ h as compared to 2.91 g/h at a feed glucose concentration of 60 g/L, but glucose conversion could not exceed that obtained with a 60 g/L feed even at a reduced flow rate of 50 mL/h. Product separation did not improve reactor performance at low glucose concentrations (120 g/L). Glucose conversion, solvent yield, and butanol selectivity decreased faster with time due to increased degeneration as compared to continuous fermentation without product removal. The specific butanol production rate of immobilized cells (ICRS) with a 60 g/L glucose feed was 0.016 g of butanol/ (g of cells-h), and butanol yield was 13.6%. This gave a specific glucose consumption rate of immobilized cells of 0.1176 g of glucose/(g of cells-h),which was 58.6% of the value for free cells. The model fits experimental results. a
Notation coefficient of temperature dependence expression of Wilson's parameter
concentration of acetic acid, g/L concentration of butanol plus butyric acid, g/L coefficient of temperature dependence expression of Wilson's parameter cell density at stage j , g of dried cell/L of reactor height of packing equivalent to an equilibrium stage vapor liquid equilibrium constant glucose saturation constant, g/L liquid molar flow rate of component i leaving stage j , mol/h total liquid molar flow rate leaving stage j , mol/h molecular weight, g; MWb, molecular weight of butanol (74 g); MW,, molecular weight of glucose (180 9); MW,, molecular weight of products total pressure at stage j , mmHg vapor pressure of component i, mmHg reaction rate, mol/(L.reactor h); rb, butanol production rate; rg, glucose consumption rate; r,, production rate of other products glucose concentration, g/L temperature, K or "C liquid molar holdup on stage j vapor molar flow rate of component i leaving stage j
total vapor molar flow rate leaving stage j , mol/h reactor volume of one stage, L liquid mole fraction of component i butanol yield vapor mole fraction of component i other product yield Greek Symbols At time interval for the relaxation method activity coefficient of component i Yi Aij Wilson's parameter specific butanol production rate, g of butanol/ (g vb of ce1ls.h) Ubmax maximum specific butanol production rate, g of butanol/ (g of ce1ls.h) w relaxation factor
Acknowledgment This work was supported by the US. Department of Energy (Contract No. DE-FG02-85-CE 40772). Literature Cited Dale, M. C.; Okos, M. R.; Wankat, P. C. An Immobilized Cell Reactor with Simultaneous Product Separation. I. Reactor Design and Analysis. Biotechnol. Bioeng. 1985a, 27,932-942. Dale, M. C.; Okos, M. R.; Wankat, P. C. An Immobilized Cell Reactor with Simultaneous Product Separation. 11. Experimental Reactor Performance. Biotechnol. Bioeng. 1985b, 27, 943-952.
Eckert, G.; Schuegerl, K. Continuous Acetone-Butanol Production with Direct Product Removal. Appl. Microbiol. Biotechnol. 1987, 27, 221-228.
Evans, P. J.; Wang, H. Enhancement of Butanol Formation by Clostridium acetobutylicum in the Presence of Decanol-Oleyl Alcohol Mixed Extractants. Appl. Enuiron. Microbiol. 1988, 54, 1662-1667.
Ferras, E.; Minier, M.; Goma, G. AcetonobutylicFermentation: Improvement of Performance by Coupling Continuous Fermentation and Ultrafiltration. Biotechnol. Bioeng. 1986,28, 523-533.
Fond, 0.;Petitdemange, E.; Petitdemange, H.; Gay, R. Effect of Glucose Flow on the Acetone Butanol Fermentation in Fed Batch Culture. Biotechnol. Lett. 1984,6, 13-18.
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Garcia, A., 111;Iannotti, E. L.; Fisher, J. L. Butanol Fermentation Liquor Production and Separation by Reverse Osmosis. Biotechnol. Bioeng. 1986, 28, 785-791. Gmehling, J.; Onken, U.; Arlt, W. Vapor-Liquid Equilibrium Data Collection; Chemistry Data Series; DECHEMA Frankfurt, Germany, 1981; Vol. 1, Parts 1, l a , 2a-d, and 5. Groot, W. J.; Schoutens, G. M.; Van Beelen, P. N.; Van den Oever, C. E.; Kossen, N. W. F. Increase of Substrate Conversion by Pervaporation in the Continuous Butanol Fermentation. Biotechnol. Lett. 1984a, 6, 789-792. Groot, W. J.; van der Oever, C. E.; Kossen, N. W. F. Pervaporation for Simultaneous Product Recovery in the Butanol/ Isopropanol Batch Fermentation. Biotechnol. Lett. 1984b,6, 709-714. Hirata, M.; Ohe, S.; Nagahama, K. Computer Aided Data Book of Vapor-Liquid Equilibria; Kodansha Limited/Elsevier Scientific Publishing Co.: Tokyo and Amsterdam, 1975. Ishii, S.; Taya, M.; Kobayashi, T. Production of Butanol by C1. acetobutylicum in Extractive Fermentation System. J.Chem. Eng. J p n . 1985, 19, 125-130. Leung, J. C. Y. Production of Acetone and Butanol by C. acetobutylicum Using Free and Immobilized Cells. Ph.D. Thesis, Massachusetts Institute of Technology, Cambridge, MA, 1982. Matsumura, M.; Kataoka, H. Separation of Dilute Aqueous Butanol and Acetone Solution by Pervaporation through Liquid Membranes. Biotechnol. Bioeng. 1987, 30, 887-895. McNeil, B.; Kristiansen, B. Effect of Temperature upon Growth Rate and Solvent Production in Batch Cultures of Clostridium acetobutylicum. Biotechnol. Lett. 1985, 7, 499-505. Monot, F.; Engasser, J. M.; Petitdemange, H. Influence of p H and Undissociated Butyric Acid on the Production of Acetone
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and Butanol in Batch Cultures of Clostridium acetobutylicum. Appl. Microbiol. Biotechnol. 1984, 19, 422-426. Moreira, A. R.; Ulmer, D. C.; Linden, J. C. Butanol Toxicity in the Butylic Fermentation. Biotechnol. Bioeng. Symp. 1981, 11, 567-579. Ounine, K.; Petitdemange, H.; Raval, G.; Gay, R. Regulation and Butanol Inhibition of D-Xylose and D-Glucose Uptake in Clostridium acetobutylicum. Appl. Enuiron. Microbiol. 1985, 49,874-878. Park, C.-H. Simultaneous Fermentation and Separation in an Immobilized Cell Trickle Bed Reactor: Acetone-Butanol-Ethanol (ABE) and Ethanol Fermentation. Ph.D. Thesis, Purdue University, West Lafayette, IN, 1989. Park, C.-H.; Okos, M. R.; Wankat, P. C. Acetone-Butanol-Ethanol (ABE) Fermentation in an Immobilized Cell Trickle Bed Reactor. Biotechnol. Bioeng. 1989, 34, 18-29. Roeffler, S. R.; Blanch, H. W.; Wilke, C. R. In-Situ Extractive Fermentation of Acetone and Butanol. Biotechnol. Bioeng. 1988, 31, 135-143. Stockhardt, J. S.; Hull, C. M. Vapor-LiquidEquilibriaand Boiling Point Composition Relations for Systems n-Butanol-Water and Isobutanol-Water. Znd. Eng. Chem. 1931,23,1438-1440. Taya, M.; Ishii, S.; Kobayashi, T. Monitoring and Control for Extractive Fermentation of C1. acetobutylicum. J. Ferment. Technol. 1985,63, 181-187. Walas, S.M. Phase Equilibria in Chemical Engineering; Butterworth Publishers: Boston, MA, 1985. Accepted January 28, 1991.
Registry No. A, 67-64-1; B, 71-36-3; E, 64-17-5.
