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Retrofit of continuous ethanol fermentation of low concentration sugar solutions by addition of a second, concentrated sugar feed
Process Biochemistry, Vol. 30, No. 4, Pp. 317-325,199s Copyright 0 I995 Elsevier Science Ltd Printed in Great Britain. All rights reserved 0032-9592/95 S9.50+0.00 0032-9592(94)00018-2
ELSEVIER
Retrofit of Continuous Ethanol Fermentation of Low Concentration Sugar Solutions by Addition of a Second, Concentrated Sugar Feed Milan pOlakovi~* & Carl-Fredrik Mandenius Department of Physics and Measurement Technology, Technology, S-581 83 Linkijping, Sweden (Received
22 December
Bioprocess
Measurement
Technology
Group, Linkaping Institute of
1993; accepted 29 April 1994)
T&e economics of ethanol production from dilute solutions may be improved by adding a concentrated sugarfeed. The production of 94% ethanolfrom a raw material containing 3% offermentable sugars, carried out in a cascade of three continuous fermenters with 75% cell recycle and in a single distillation column, was taken as a reference case. The parameter of variation was the ratio of the concentrated feedflow containing 50% of sugar to the dilute feed flow. The effects of the additional sugar feed on the fermentation, centrifugation and distillation capacities of the plant were analysed and the retrofit of the fermentation unit was optimized using the fermenter volume and ceil recycle as optimization parameters. Approximate economic evaluations of retrofit were made by comparing some essential costs in the retrofitted system to these costs in the above rqference system and in a conventionalfermentation plant, respectively.
NOTATION Concentration vector (kg/m3) Ethanol concentration (kg/m3) Substrate concentration (kg/m3) Cell concentration on the dry mass basis (kg/m3) Centrifuge capacity (m3/h) Distillate mass flow (kg/h) Function Distillation feed (kg/h) Suspension flow rate to centrifuge (m3/h) Product inhibition constant (97-9 kg/m3) *To whom correspondence should be addressed. Present address: Department of Chemical and Biochemical Engineering, Slovak %chnical University, Radlinskeho 9, 8 1237 Bratislqva, Slovakia.
MR P. 2 ‘P rb
2
RR VF vrn XF
317
Monod constant (2.0 kg/m3) Mass flow of dilute sugar feed (kg/h) Mass flow of concentrated sugar feed (kg/h) Mixing ratio Parameter vector Condenser duty (W ) Reboiler duty (W ) Ethanol production rate (kg/m3/h) Substrate consumption rate (kg/m3/h) Cell growth rate (kg/m3/h) Reflux ratio Cell recycle ratio Fermenter volume (m’) Bottom vapour flow (kmol/h) Mass fraction of ethanol in the distillation feed
318
Milan PolakoviE; Carl-Fredrik Mandenius
YXS
Ethanol yield (0.46) Cell yield (0.099)
71 Pin
Centrifuge efficiency Maximum specific growth rate (0.48 h-l)
Yps
Subscripts Conventional fermentation C R Retrofitted case 0 Reference case
INTRODUCTION Since the oil crisis of 1973, the biological production of ethanol has become one of the top candidates for the replacement of petrochemicals for energy usage in addition to use as a chemical feedstock. Numerous studies have been performed dealing with various aspects of ethanol fermentation. Technologically feasible solutions do not however guarantee a breakthrough in the usage of renewable resources unless the prices of products obtained are competitive with those from fossil raw materials. For that reason, the economics of ethanol fermentation has been investigated more than those of any other biotechnological process. The first studies dealing with the process economics clearly concluded that ethanol fermentation was not yet a competitive alternative but that its potential should nevertheless not be dismissed.‘y2 An extensive economic evaluation of ethanol fermentation, including different fermenter configurations such as batch, continuous tank (simple and series), tower, and immobilized cell fermenters, has been made by Maiorella et aL3 Similar issues have been tackled in a recent study using more current economic data.4 Several studies have dealt with the economic assessment of using cellulose hydrolysate, either from waste1-3g5 or wood.6 One disadvantage with the application of these materials is their low sugar content resulting in low cell and ethanol concentrations. Since the distillation cost per unit amount of ethanol produced is substantially higher at low ethanol concentrations,7 several investigators have dealt with the idea of concentrating sugar solutions prior to fermentation.‘,“*7 Contrary to earlier studies,‘a3 Zacchi and Axelsson questioned evaporation as a favourable option but claimed economic feasibility of reverse osmosis.’
The main economic objective of cell recycling is to decrease the fermenter size. The economics of cell recycling has been the subject of several studies.3,8,y Maiorella et al. considered centrifugation, which is currently the most used separation method for cell recycling in industry, and suggested a recycle ratio where 90% of all cells in the fermenter were recycled.3 A corresponding fermenter cell concentration of 100 kg/m3 is at the upper limit achievable with this technique. Continuous ethanol plants in industry operate at substantially lower cell concentrations. Membrane filtration, a technique not yet applied on a large scale, was evaluated by Groot et al.” in a high density fermentation with a cell concentration of 150 kg/m’. Following a refined kinetic modellO and using membrane filtration for cell recycling, Warren et aLa found an optimum value of cell concentration at 54 kg/m3. This paper presents another alternative for improving the economics of continuous plant processing dilute sugar solutions: adding a second substrate feed with higher sugar content. A reference process was retrofitted using different ratios of substrate feed flows. For each retrofitted case, the changes required in the plant design and operation are evaluated separately for fermentation and distillation. The economic effects of retrofit are also estimated.
MATHEMATICAL SIMULATION
MODELS
AND
Although we tried to make the study as general as possible, the number of parameters studied had to be limited. For that reason, the value of the sugar concentration in the dilute feed was chosen to be 30 kg/m3, which is a typical value for spent sulphite liquor fermentation. This information, as well as some other facts employed in this study, were obtained from an ethanol plant in the North of Sweden. The concentrated sugar feed contained 50 mass % of sugar. Fermentation kinetics The kinetics of ethanol fermentation on spent sulphite liquor has recently been investigated.11,12 Unfortunately, these authors did not remove the sulphur dioxide from the spent sulphite liquor, so their kinetic equations give very low fermentation rates. Mohrmann and Wenzeli3 showed that after the complete removal of SO,, the residence time
Retrofit of ethanolfermentation of low concentrationsugar solutions
in a continuous fermenter could be decreased to 2.5 h, which essentially does not differ from the maximum residence times attainable in a conventional fermentation. This result agrees with our experience from large scale plant measurements.14 Since the cell and ethanol concentrations in this study were relatively low, we employed the Ghose and Tyagi model15 with the kinetic parameters determined by Hill and Robinson.16 This model follows the Monod-type limitation of cell growth on sugar substrate and includes an inhibition term which predicts a linear decline in the growth rate with the ethanol concentration:
where r, is the cell growth rate expressed as the amount of dry mass produced per unit volume and time, c, is the cell concentration, c, is the substrate concentration, cp is the ethanol concentration, ,u,,, is the maximum specific growth rate, K, is the Monod constant and K, is the product inhibition constant. In this model the substrate consumption rate, r,, and ethanol production rate, rp, are related to the cell growth rate, r,, with simple linear relations assuming constant yield factors. The values of the kinetic parameters are given in the Notation. It is important to emphasize here that all results presented are independent of the value of the maximum specific growth rate, ,u,,,. The fermenter material balance can be rearranged into the following form: (2)
~L%=f,(c?P)
319
where V, is the fermenter volume, fi is a function of the concentration vector, c, and the parameter vector, p, including the flow rates of single streams and remaining kinetic parameters. Obviously, if two cases are compared, the following relation is valid:
~fhIl)
+ F2
Fermentation unit Throughout this study the flowsheet configurations of the fermentation unit differed but a typical layout was identical with the one presented in Fig. 1. It consisted of three fermenters in series, a continuous centrifuge, and a flow splitter. The first fermenter was fed with two sugar streams and the cell recycle stream. Steady-state simulations of the fermentation unit were made using the equation-oriented simulator SPEEDUP (Aspen Technology, Inc., Cambridge, MA, USA). Ideal mixing in the fermenter was assumed. As there are almost no data available for the process simulation of centrifuges, a simplified model was suggested. No cell loss in the effluent was assumed and a cell concentration in the slurry outflow was set at a conservative estimate of 1.50 kg/m3. Due to the cell recycle, all process units had to be solved simultaneously. A constant value of the dilute sugar feed flow was used in all cases. In the reference case, no concentrated sugar substrate was fed into the fermenter cascade. The cell recycle ratio, RR, was defined as the ratio of flow fed back to the fermenter to the total slurry flow from the centrifuge. The recycle ratio was set to 0.75 in the refer-
Concentratedfeed 0
Centdfuge Distillation
Dilute feed G Fermentor
I Cell recycle
b
Cell bleed
Fig. 1.
Schematic flowsheet of fermentation
unit.
320
Milan Polakovic’, Carl-Fredrik Mandenius
ence case but was calculated in the retrofit design. The flow of the concentrated sugar feed in the retrofits was defined by a chosen value of the mixing ratio, MR:
MR=m, m,
(4)
where &, and A, are the mass flows of the dilute and concentrated sugar feeds, respectively. The residual sugar concentration at the outlet of the fermenter cascade was 0.5 kg/m3. All fermenters in the cascade had the same volume. All results were expressed per unit mass of 94% ethanol. Distillation unit A significant consideration here was the selection of a distillation set-up. There has been a strong interest in energy saving alternatives for ethanol recovery in the recent past and these techniques are utilized in fermentation plants.17 However, their computations are more time consuming than that of conventional single column distillation. The results of a recent study showed that the ratio of costs for the distillation of two feeds of different ethanol concentration (2.5 versus 4.5% ethanol) was practically independent of the distillation set-up. I8 Thus, the use of a single column was satisfactory for the purpose of this study. The column had 40 theoretical stages and the pressure was considered to be atmospheric and constant along the column. The calculations were performed with the ASPEN PLUS simulator (Aspen Technology, Inc., Cambridge, MA, USA) using the built-in rigorous column model. The Wilson equation was used for the evaluation of activity coefficients in the liquid phase and the Redlich-Kwong state equation for the vapour phase. The feed temperature was set at 30°C. Two prerequisites for the distillation design were 94 mass % of ethanol in the distillate and 99% recovery of ethanol from the feed.
RETROFIT
OF FERMENTATION
UNIT
Figure 2 shows the effect of the retrofit on the production capacity of the fermentation unit. The plant capacity can be doubled by adding the concentrated sugar feed in a mixing ratio of O-06. The maximum capacity in the range of mixing ratios investigated is almost eight times higher than in the reference case. Simultaneously, the output
0.2 0.3 Mixing ratio (-)
0.1
Fig. 2. Effect of retrofit on the production capacity of fermentation unit. Production capacity is given as a relative value of the production capacity of the reference case when the mixing ratio was equal to 0. The second curve represents the outlet ethanol concentration from the fermentation unit.
I6I
0.6
-
0
0.2
0.4
0.6
0.8
1
Split ratio C-1
sugar feed Fig. 3. Effect of distributing the concentrated on the volume of the cascade without cell recycle for a mixing ratio of 0.2. The split ratio represents the ratio of flow into the second stage to the total flow of concentrated feed. The fermenter volume is given in a relative form when it is compared to the volume at a split ratio equal to 0.
ethanol concentration from the fermentation unit increased from 1.3 to 7.5%. The results shown in this figure are not dependent on the fermentation unit set-up, but only on the stoichiometry and sugar conversion. In some previous studies dealing with multistage continuous ethanol fermentation, molasses was fed into several stages.‘9,20 In this case, the effect of distributing the concentrated sugar feed into the first and second stages on the retrofit design was studied. The last fermenter in the cascade could not be fed since this must operate at a low substrate concentration. It was found that feeding the second stage either did not affect the retrofit or was disadvantageous. Figure 3 illustrates an example where up to 60% of concentrated sugar feed could be fed to the second stage
Retrofit of ethanolfermentation
almost without any changes in the fermenter volume. This conclusion was only valid for mixing ratios smaller than 0.2. Above this value, the maximum fraction of the concentrated feed which could be fed into the second fermenter rapidly decreased. Consequently, multi-feeding was not considered in further calculations. As mentioned above, adding the concentrated sugar feed can increase the production capacity but this also means a higher flow rate. On the other hand, the higher amount of sugar fed into the system would require a longer residence time. Figure 4 illustrates that in some cases this demand might be satisfied without the installation of an additional fermenter. For example, only a 7% larger fermenter volume is required for doubling the production capacity in the system without recycle as follows from Figs 2 and 4. Besides the fermenter capacity, the centrifuge capacity must bc considered in the retrofit. The centrifuge capacity is determined by the suspension flow rate as well as by the suspension concentration. A higher concentration of solids implies a lower suspension flow rate that can be put through the centrifuge. Since in our case, adding the concentrated sugar feed means both higher flow rate and cell concentration, an increase in the centrifuge capacity is inevitable. The capacity of a centrifuge is generally a function of the solid size distribution and of the solid and liquid densities. In this study, the centrifuge capacity, CC, was then constant and represented a limit value of suspension flow rate for the yeast cell concentration close to 0. For the relation of the centrifuge capacity to the suspension flow rate, F,, we defined a quantity, called the centrifuge efficiency, r :
of low concentration
321
sugar solutions
According to vendor documentation on the commercial large scale centrifuges, with the capacity in the range of 75-200 m3/h of yeast cell suspension, the centrifuge efficiency is primarily dependent on the cell concentration but has very little dependence on the centrifuge size.*l The adaptation of these data gave the following polynomial relation of the centrifuge efficiency on the cell concentration (Fig. 5): ~=1-2~35x10-3c,-2-84x10-‘c2
X
(6)
Figure 5 shows that at a cell concentration of 100 kg/m3 the suspension flow rate amounts only to 50% of the centrifuge capacity. Figure 6 shows the relationship between the centrifuge capacity and the addition of concentrated feed assuming that the fermenter volume is the same as in the reference system. For a low mixing ratio, O-1, the relative increase in the centrifuge capacity is practically constant below a recycle ratio of 0%. This increase roughly corresponds to the flow of concentrated feed. On the other hand the required increase of centrifuge capacity is very high at a mixing ratio of O-4. As Fig. 6 shows, if the reference system has a recycle ratio above 0.4, which is relevant to continuous ethanol fermentations in industry, the system cannot be retrofitted by increasing the centrifugation capacity alone since this would require unrealistically large centrifugation capacities. The results above imply that a simultaneous increase in the fermenter volume and centrifuge capacity is required in some cases to retrofit the system, particularly when the concentrated sugar
2t 5
‘5 ;F
0.6
0.6
T
1
t_,,. 00
0.1
,
,
0.2
0.3
1 0.4
Mixing ratio (-)
Effect of the mixing ratio on the total volume of the Fig. 4. cascade without cell recycle. The fermenter volume is given as a relative value of the volume at a mixing ratio equal to 0 (reference case).
3 g z
0.4 0.2
u
0
Centrifuge Fig. 5. concentration.
efficiency
as a function
of yeast
cell
322
Milan Polakovi?,
Carl-Fredrik
feed ought to be added at a higher proportion. It was assumed that the total fermenter volume could be changed only by adding one or more fermenters of equal volume. The decision as to whether an extra fermenter should be used or not was based on a comparison of the costs of centrifugation and fermentation. We used an approximation that the costs of the installed fermenter to
centrifuge capacity were 3 : 1. This means that doubling the centrifuge capacity was economically equivalent to adding a fourth fermenter. Table 1 presents the retrofit results for all mixing ratios studied. It is shown here that a multiple increase in production capacity could be achieved with much less than a proportional increase in ferementer and centrifuge capacity. No additional fermenter installation is suggested at a mixing ratio of 0.1, which corresponds to a 150% increase in production capacity. The last column in Table 1 gives a multiplication factor characterizing the required increase in the centrifuge capacity of the retrofitted system. This factor varied between 37 and 156%. A certain increase in the centrifuge capacity might be managed without additional centrifuge installation depending on how far below the maximum capacity the centrifuge operates. Centrifuges, with the capacity relevant to the industrial ethanol fermentation, are manufactured in discrete sizes with a scale factor of 2~5.~~ This can give a significant difference between designed and utilized centrifuge capacity and can also give a space for the increased capacity required at retrofit. The economic evaluation of the retrofit of the fermentation unit followed from the results presented in Table 1. Since the retrofit was made with fermenters of equal volume, a linear approximation between the fermentation cost (excluding substrates) and total fermenter volume was appropriate. The centrifugation cost relative to the centrifuge capacity can also be expressed through a linear relationship.21’22 Thus, the comparisons of fermenter volumes and centrifuge capacities were also comparisons of the costs of the corresponding operations. Columns 3-6 of Table 2 compare the costs of fermentation and centrifugation with the refer-
r
0
02
0.4
0.8
RR,
0.8
1
t-1
Fig. 6. Relation between the centrifuge capacity and the mixing ratio in a cascade with recycle, expressed through the dependence of the relative centrifuge capacity (retrofitted versus reference system) on the cell recycle ratio in the referencc system with the mixing ratio as curve parameter: 1, mixing ratio 0.1; 2, mixing ratio 0.4. The dashed parts of the curves indicate the region where the recycle ratio in the retrofitted system is above the chosen, practical limit value of 0.9.
Table 1. Results of the retrofit of the fermentation unit (RR,, cell recycle ratio; c,;, cell concentration; MA, mixing ratio; V,,, fermenter cascade volume; CC,, centrifugation capacity). The subscripts o and R denote the reference and retrofitted cases, respectively RR, (CA
MR
0.75 (11.1)
0.1 0.2 0.3 0.4
V,,lv,
cxR
1 1.33 1567 2.33
32.6 43.3 56.2 61.0
CC&C”
RR,
082 0.72 0.83 082
Mandenius
1.37 1.70 2.22 2.56
Table 2. Cost related results of the retrofit of the fermentation unit (D,, mass tlow of ethanol obtained in the distillation unit). The subscript C denotes the conventional fermentation case. For the explanation of other symbols see Table 1 RR,
0.75
MR
v,
v,,
- vFu
D A
D,-0,
!!Ea
AV
D,
Do
cc, D _.A cc, D”
cc,
- cc,
D, -Do
VFR D A
cc,
v,
- v,,
cc,
- cc,
DR - Do
cc, D A
I.fx
EL
cc,
D*
DC
DC
DC
DC
0
4.95 2.52
1.08
o-12 0.16 0.24
1.92 1.81 1.63
1.02 1.19 l-14
0 0.1
0.37
0
0.51
0.22
1.23 0.46
;:; 0.4
0.3 O-27 0.3
0.13 0.10 0.20
0.39 0.37 0.33
0.2 1 0.24 0.23
0.37 0.34 0.37
D,-D, cc,
Retrofit of ethanol fermentation
ence case. All results are expressed per unit mass of 94% ethanol produced. While the odd columns show the relative costs for the retrofitted unit, the even columns represent the estimates of expenses per unit mass of extra ethanol obtained after the retrofit. The unit cost for both fermentation and centrifugation showed a significant decrease with the mixing ratio to 50% and less compared with the reference case. Both expenses decreased further with increasing mixing ratio. In general, the centrifugation cost varied over a smaller range than the fermentation cost. Even more convincing results were obtained after separating the cost of the reference system. In one case the fermentation expenses for additionally gained ethanol were equal to zero and their maximum value was 20% of the expenses in the reference system. The additional centrifugation cost varied only slightly and was somewhat more than 20% of that in the reference system. These figures demonstrate clearly the potential of the retrofit under study. Due to the decisive role of the substrate cost, it was of interest to compare the economics of the retrofitted plant with a conventional fermentation using substrates with high sugar content. This comparison is shown in the last four columns in Table 2. As a reference conventional fermentation was taken a system consisting of three fermenters in series fed with a substrate with a concentration of 164.3 kg/m3, and with a cell recycle of OS. The chosen substrate concentration corresponds to a total amount of sugar fed in the case of the mixing ratio of 0.4. The cell concentration at the outlet of the fermenter cascade was then 29.5 kg/m3. The advantages of the retrofit compared to the conventional fermentation were emphasized when the net expenses on additionally gained ethanol were evaluated. Although the extra centrifugation cost was about the same, the extra fer-
of low concentration
sugar solutions
323
mentation cost showed nearly the same positive outcome as in the previous comparison to the reference case (see columns 4 and 8 in Table 2).
RETROFIT
OF DISTILLATION
UNIT
Table 3 reviews the results of the distillation column calculations for different feed concentrations, XF, which correspond to the four different mixing ratios studied in the retrofit of the fermentation unit. As mentioned above, the amount of ethanol produced increased substantially with the mixing ratio, almost eight times compared to the reference case. As the distillate flow in none of these cases was above 10% of the feed flow value, the relative increase in the vapour flow at the column bottom was substantially smaller than the increase in the amount of distillate. The ratio of the bottom vapour flows was practically identical to the ratio of the reboiler duties. The relatively low consumption of heat in the reboiler at higher mixing ratios was also partially due to the decreased reflux ratio. The cooling duty of the condenser increased almost as sharply as the amount of distillate but was always smaller than the reboiler duty. The above results indicated clearly that adding the concentrated sugar feed would require the expansion of distillation capacities. The 30% increase in vapour flow at a mixing ratio of 0.1 could possibly be accommodated in the distillation column designed for the dilute ethanol feed but in any case the condenser capacity had to be increased significantly. The results presented in Table 3 were used for the cost evaluation of the distillation retrofit. Since the distillation column forms by far the largest portion of the total capital cost and the relative changes in the costs of the distillation column and reboiler were considered to be approximately
Table 3. Results of the retrofit of the distillation unit (MR, mixing ratio; &mass fraction of ethanol in the distillation feed; D/F, distillate to feed mass flow ratio; R, reflux ratio; D,, distillate flow; Q,,,,, Q,,,, reboiler and condenser duties, respectively; Vmi, bottom vapour flow). The subscripts o and R denote the reference and retrofitted cases, respectively
o-o
0.1 0.2 0.3 o-4
1.36 3.32 4.96 6.35 7.53
1.43 3.50 5.22 6.68 7.93
1 2.69 4.39 6.08 7.77
3.28 2-69 2.5 2.42 2.39
1 1.29 1.57 1.85 2.14
i.32 3.58 4.86 6.16
0.163 0.296 0.372 0.427 @469
324
Milan Polakovit, Carl-Fredrik Mandenius
Table 4. Cost-related
For the explanation MR
results of the retrofit of the distillation unit. The subscript C denotes the conventional of uther symbols see Table 3
e,,L
CL,< - dw,,
DtZD, L
v,,,
Do 0 0.1 0.2 o-3 0.4
K,,H - v,,,
D,-Do-D,-D,,
Do
Q,IIHV _I?1L!
Q,,v,,,,
Do
Do
D,
-
1.30
0.16 0.16 0.16
equal, a simplified but reasonably accurate cost evaluation was made. A linear approximation between the column cost including reboiler and vapour flow rate was considered. This approximation was compared with the empirical equation derived by Zacchi and Axelsson.7 In the range of the flow rates used, the deviation from the correct value was at a maximum at 10%. This deviation was compensated for by the condenser cost, which was otherwise neglected. The reboiler steam consumption was considered to be a key component of distillation operation cost. As this was proportional to the reboiler duty, a constant proportion between the distillation capital and operation costs was obtained in agreement with earlier results.’ Similarly as in the case of the fermentation unit, the economic consequences of retrofit were compared with the reference case as well as with conventional fermentation. The second and third columns of Table 4 arc rclatcd to the former process, while the last two columns of Table 4 to the latter. The second column of Table 4 shows that the estimated distillation cost per unit mass of final ethanol product already dropped to 48% at a mixing ratio of 0.1 and went further down to 28% at the mixing ratio of 0.4. The third column of Table 4 shows that the cost of retrofit per unit mass of additionally gained ethanol was independent of the mixing ratio and was only 16% of the cost in the reference plant. Obviously, the distillation cost per unit mass of ethanol (fifth column of Table 4) was higher than in the conventional fermentation. But, after the cost for the reference plant had been subtracted, the retrofit cost (last column of Table 4) showed a positive balance. The cost per unit mass of additionally gained ethanol showed a constant value of 60% of the cost in the conventional fermentation.
1.10
1
DISCUSSION
v,,
- v,,
i)h’(L D,
3.64 1.I4
0.16 0.36 0.30 0.28
- a,
case.
D,-D,,-D,-D,
n,-n,
e,L
ad
fermentation
D,
D,
0.60 0.60 0.60 0.61
AND CONCLUSION
The key factor in the evaluation of the economics of double sugar feed fermentation is the ratio of substrate costs. Two principal options can be defined here. First, that the costs of both substrates on the total mass basis are comparative and secondly, that the low content sugar substrate is substantially cheaper. Both alternatives are reasonable. The cost of cellulose hydrolysate, even from waste materials, is rather high due to the high cost of conversion of cellulose to fermentable sugars.” On the other hand, some wastes containing fermentable sugars like spent sulphite liquor or whey, arc rather cheap, especially when they are further utilized on the production site. For example, the calculated price of spent sulphite liquor in the pulp plant is 0.8 USS per m3 of liquor. This low price of the liquor makes the production of technical ethanol profitable in spite of the low sugar content. The price of molasses is currently about 325 US$ per tonne in Sweden and the transportation cost amounts to 75 US$ per tonne resulting in a total cost of approximately 200 US$ per tonne of molasses. The economic effects of the retrofit presented in Tables 2 and 4 are related to the above two substrate cost ratios in a different way. For cellulose hydrolysate, the bcncfits of the retrofit are better demonstrated if the costs are compared to the reference process. The retrofit improves the economy of the whole process and brings down the cost of ethanol produced in this way. For a cheap substrate, it is more relevant to compare the costs with conventional fermentation since, after adding a concentrated sugar feed, the total cost of produced ethanol will increase but this will be economically justified while the ethanol cost remains lower than the cost achieved in a conventional fermentation plant.
Retrofit of ethanol fermentation of low concentration
A relatively small increase in the mixing ratio of sugar feeds resulted in a substantial increase in the plant production capacity which could be accommodated in the existing plant facilities with minor modifications. A plant designed for a certain ratio of sugar feeds would thus allow rather flexible operation to accommodate seasonal fluctuations in feedstock amounts or price fluctuations.
ACKNOWLEDGEMENTS This work was supported by a grant from the Swedish Ethanol Foundation and the National Board for Technical Development (NUTEK). Professor Guido Zacchi is kindly acknowledged for his advice on ethanol distillation. We are grateful to Mr Hans Axelsson, from Alfa Lava1 Separation AB for the centrifuge documentation and personal information. We thank MS Eva Larsson, from MoDo Paper AB, t)rnskiildsvik and MS Tngrid Ascard, from Sockerbolaget AB, Malmii for providing information on the costs of spent sulphite liquor and molasses, respectively.
7.
8. 9.
10.
1 I.
12.
13. 14.
15.
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sugar solutions
325
ethanol production for SSF technology. Appl. Biochem. Biotechnol., 34/35 (1992) 639-49. Zacchi. G. & Axelsson. A.. Economic evaluation of meconcentration in product& of ethanol from dilute sugar solutions. Riotechnol. Bioengng, 34 (1989) 223-33. Warren, R. K., Hill, G. A. & Macdonald, D. G., Simulation and optimization of a cell recycle system. Appl. Biothem. Bioierhnol., 34135 (1992) 585-45. __ Groot. W. J.. Kraavebrink. M. R., van der Lans, R. G. J. M. RLLuyben, K. -Ch. A. M., Ethanol production in an integrated fermentation/membrane system. Process simulations and economics. Bioproc. Engng, 8 (1993) 189-20 I. Warren, R. K., Hill, Cr. A. & Macdonald, D. G., Imoroved bioreaction kinetics for the simulation of continuous ethanol fermentation by Saccharomyces cerevisiae. Biotechnol. from., 6 (1990) 3 19-25. Safi, B. F., Rouleau, D., Mayer, R. C. & Desrochers, M., Fermentation kinetics of spent sulfite liquor by Saccharomyces cerevkiae. Riotechnol. Bioengng, 28 (1986) 944-51. Rvusseau, S., Rouleau, D., Yerushalmi, L. & Mayer, R. C., Effect of temperature on fermentation kinetics of waste sulfite liauor bv Sxcharomvces cerevisiae. J. Chem. Technol. biotech;ol., 53 (1995) 285-91. Mohramann. K. H. & Wenzel. W. D., Effect of residence time at spent sulfite liquor fermentation. Chem. Techn&.,22(1992)201-3(inGerman). Polakovi& M., Hiirnsten, E. G. & Mandenius, C. F., Simulation of sDent sulfite liquor fermentation using the object oriented knowledge -based shell G2. &ma1 LiTH-IFM-R-199, Link&in2 revort . _ University, Linkiiping, Sweden, 1992. Ghose, T. K. & Tyagi, R. D., Rapid ethanol fermentation of cellulose hydrolyzate. II. Product and substrate inhibition and optimization of fermenter design. Biotechnol. Bioengng, 21 (1979) 1401-20. Hill, G. A. & Robinson, C. W., A modified Ghose model for batch cultures of Saccharomyces rerevkiae at high ethanol concentrations. Chem. Engng J. Biochem. Engng J., 44 i 1990) B69-80. Serra, A., Poch, M. & Sol& C., A survey of separation svstems for fermentation ethanol recovcrv. !‘rocess Lkochem., 22 (I 987) 154-K Larsson. M. & Zacchi. G.. Energv saving alternatives within distillation. Internal reporF LUTgDH/(TKKA7001 )/l-32/( 1993), Lund Institute of Technology, Lund, Sweden, 1993 (in Swedish). Chattaway, T., Goma, G. & Renaud, P.-Y., Modelling ethanol and secondary inhibitions of ethanol fermentation in a multistage reactor. Biotechnol. Rioengng, 32 (1988) 271-6. Chen, H. C. & Mou, D. G., Pilot-scale multi-stage multifeeding continuous ethanol fermentation using nonsterile cane molasses. Biotechnol. Lett.. 12 (1990) 367-72. Alja-L.aval Separation Equipment Documentation. AlfaLava1 Sevaration AB, Tumba. Sweden. 1993. Walas, S..M.. Costs individual equipment. In Chemical Process Equipment. Selection and Design. Butterworths, Boston, MA, 1988, Chapter 20, pp. 663-9.
of
ELSEVIER
Retrofit of Continuous Ethanol Fermentation of Low Concentration Sugar Solutions by Addition of a Second, Concentrated Sugar Feed Milan pOlakovi~* & Carl-Fredrik Mandenius Department of Physics and Measurement Technology, Technology, S-581 83 Linkijping, Sweden (Received
22 December
Bioprocess
Measurement
Technology
Group, Linkaping Institute of
1993; accepted 29 April 1994)
T&e economics of ethanol production from dilute solutions may be improved by adding a concentrated sugarfeed. The production of 94% ethanolfrom a raw material containing 3% offermentable sugars, carried out in a cascade of three continuous fermenters with 75% cell recycle and in a single distillation column, was taken as a reference case. The parameter of variation was the ratio of the concentrated feedflow containing 50% of sugar to the dilute feed flow. The effects of the additional sugar feed on the fermentation, centrifugation and distillation capacities of the plant were analysed and the retrofit of the fermentation unit was optimized using the fermenter volume and ceil recycle as optimization parameters. Approximate economic evaluations of retrofit were made by comparing some essential costs in the retrofitted system to these costs in the above rqference system and in a conventionalfermentation plant, respectively.
NOTATION Concentration vector (kg/m3) Ethanol concentration (kg/m3) Substrate concentration (kg/m3) Cell concentration on the dry mass basis (kg/m3) Centrifuge capacity (m3/h) Distillate mass flow (kg/h) Function Distillation feed (kg/h) Suspension flow rate to centrifuge (m3/h) Product inhibition constant (97-9 kg/m3) *To whom correspondence should be addressed. Present address: Department of Chemical and Biochemical Engineering, Slovak %chnical University, Radlinskeho 9, 8 1237 Bratislqva, Slovakia.
MR P. 2 ‘P rb
2
RR VF vrn XF
317
Monod constant (2.0 kg/m3) Mass flow of dilute sugar feed (kg/h) Mass flow of concentrated sugar feed (kg/h) Mixing ratio Parameter vector Condenser duty (W ) Reboiler duty (W ) Ethanol production rate (kg/m3/h) Substrate consumption rate (kg/m3/h) Cell growth rate (kg/m3/h) Reflux ratio Cell recycle ratio Fermenter volume (m’) Bottom vapour flow (kmol/h) Mass fraction of ethanol in the distillation feed
318
Milan PolakoviE; Carl-Fredrik Mandenius
YXS
Ethanol yield (0.46) Cell yield (0.099)
71 Pin
Centrifuge efficiency Maximum specific growth rate (0.48 h-l)
Yps
Subscripts Conventional fermentation C R Retrofitted case 0 Reference case
INTRODUCTION Since the oil crisis of 1973, the biological production of ethanol has become one of the top candidates for the replacement of petrochemicals for energy usage in addition to use as a chemical feedstock. Numerous studies have been performed dealing with various aspects of ethanol fermentation. Technologically feasible solutions do not however guarantee a breakthrough in the usage of renewable resources unless the prices of products obtained are competitive with those from fossil raw materials. For that reason, the economics of ethanol fermentation has been investigated more than those of any other biotechnological process. The first studies dealing with the process economics clearly concluded that ethanol fermentation was not yet a competitive alternative but that its potential should nevertheless not be dismissed.‘y2 An extensive economic evaluation of ethanol fermentation, including different fermenter configurations such as batch, continuous tank (simple and series), tower, and immobilized cell fermenters, has been made by Maiorella et aL3 Similar issues have been tackled in a recent study using more current economic data.4 Several studies have dealt with the economic assessment of using cellulose hydrolysate, either from waste1-3g5 or wood.6 One disadvantage with the application of these materials is their low sugar content resulting in low cell and ethanol concentrations. Since the distillation cost per unit amount of ethanol produced is substantially higher at low ethanol concentrations,7 several investigators have dealt with the idea of concentrating sugar solutions prior to fermentation.‘,“*7 Contrary to earlier studies,‘a3 Zacchi and Axelsson questioned evaporation as a favourable option but claimed economic feasibility of reverse osmosis.’
The main economic objective of cell recycling is to decrease the fermenter size. The economics of cell recycling has been the subject of several studies.3,8,y Maiorella et al. considered centrifugation, which is currently the most used separation method for cell recycling in industry, and suggested a recycle ratio where 90% of all cells in the fermenter were recycled.3 A corresponding fermenter cell concentration of 100 kg/m3 is at the upper limit achievable with this technique. Continuous ethanol plants in industry operate at substantially lower cell concentrations. Membrane filtration, a technique not yet applied on a large scale, was evaluated by Groot et al.” in a high density fermentation with a cell concentration of 150 kg/m’. Following a refined kinetic modellO and using membrane filtration for cell recycling, Warren et aLa found an optimum value of cell concentration at 54 kg/m3. This paper presents another alternative for improving the economics of continuous plant processing dilute sugar solutions: adding a second substrate feed with higher sugar content. A reference process was retrofitted using different ratios of substrate feed flows. For each retrofitted case, the changes required in the plant design and operation are evaluated separately for fermentation and distillation. The economic effects of retrofit are also estimated.
MATHEMATICAL SIMULATION
MODELS
AND
Although we tried to make the study as general as possible, the number of parameters studied had to be limited. For that reason, the value of the sugar concentration in the dilute feed was chosen to be 30 kg/m3, which is a typical value for spent sulphite liquor fermentation. This information, as well as some other facts employed in this study, were obtained from an ethanol plant in the North of Sweden. The concentrated sugar feed contained 50 mass % of sugar. Fermentation kinetics The kinetics of ethanol fermentation on spent sulphite liquor has recently been investigated.11,12 Unfortunately, these authors did not remove the sulphur dioxide from the spent sulphite liquor, so their kinetic equations give very low fermentation rates. Mohrmann and Wenzeli3 showed that after the complete removal of SO,, the residence time
Retrofit of ethanolfermentation of low concentrationsugar solutions
in a continuous fermenter could be decreased to 2.5 h, which essentially does not differ from the maximum residence times attainable in a conventional fermentation. This result agrees with our experience from large scale plant measurements.14 Since the cell and ethanol concentrations in this study were relatively low, we employed the Ghose and Tyagi model15 with the kinetic parameters determined by Hill and Robinson.16 This model follows the Monod-type limitation of cell growth on sugar substrate and includes an inhibition term which predicts a linear decline in the growth rate with the ethanol concentration:
where r, is the cell growth rate expressed as the amount of dry mass produced per unit volume and time, c, is the cell concentration, c, is the substrate concentration, cp is the ethanol concentration, ,u,,, is the maximum specific growth rate, K, is the Monod constant and K, is the product inhibition constant. In this model the substrate consumption rate, r,, and ethanol production rate, rp, are related to the cell growth rate, r,, with simple linear relations assuming constant yield factors. The values of the kinetic parameters are given in the Notation. It is important to emphasize here that all results presented are independent of the value of the maximum specific growth rate, ,u,,,. The fermenter material balance can be rearranged into the following form: (2)
~L%=f,(c?P)
319
where V, is the fermenter volume, fi is a function of the concentration vector, c, and the parameter vector, p, including the flow rates of single streams and remaining kinetic parameters. Obviously, if two cases are compared, the following relation is valid:
~fhIl)
+ F2
Fermentation unit Throughout this study the flowsheet configurations of the fermentation unit differed but a typical layout was identical with the one presented in Fig. 1. It consisted of three fermenters in series, a continuous centrifuge, and a flow splitter. The first fermenter was fed with two sugar streams and the cell recycle stream. Steady-state simulations of the fermentation unit were made using the equation-oriented simulator SPEEDUP (Aspen Technology, Inc., Cambridge, MA, USA). Ideal mixing in the fermenter was assumed. As there are almost no data available for the process simulation of centrifuges, a simplified model was suggested. No cell loss in the effluent was assumed and a cell concentration in the slurry outflow was set at a conservative estimate of 1.50 kg/m3. Due to the cell recycle, all process units had to be solved simultaneously. A constant value of the dilute sugar feed flow was used in all cases. In the reference case, no concentrated sugar substrate was fed into the fermenter cascade. The cell recycle ratio, RR, was defined as the ratio of flow fed back to the fermenter to the total slurry flow from the centrifuge. The recycle ratio was set to 0.75 in the refer-
Concentratedfeed 0
Centdfuge Distillation
Dilute feed G Fermentor
I Cell recycle
b
Cell bleed
Fig. 1.
Schematic flowsheet of fermentation
unit.
320
Milan Polakovic’, Carl-Fredrik Mandenius
ence case but was calculated in the retrofit design. The flow of the concentrated sugar feed in the retrofits was defined by a chosen value of the mixing ratio, MR:
MR=m, m,
(4)
where &, and A, are the mass flows of the dilute and concentrated sugar feeds, respectively. The residual sugar concentration at the outlet of the fermenter cascade was 0.5 kg/m3. All fermenters in the cascade had the same volume. All results were expressed per unit mass of 94% ethanol. Distillation unit A significant consideration here was the selection of a distillation set-up. There has been a strong interest in energy saving alternatives for ethanol recovery in the recent past and these techniques are utilized in fermentation plants.17 However, their computations are more time consuming than that of conventional single column distillation. The results of a recent study showed that the ratio of costs for the distillation of two feeds of different ethanol concentration (2.5 versus 4.5% ethanol) was practically independent of the distillation set-up. I8 Thus, the use of a single column was satisfactory for the purpose of this study. The column had 40 theoretical stages and the pressure was considered to be atmospheric and constant along the column. The calculations were performed with the ASPEN PLUS simulator (Aspen Technology, Inc., Cambridge, MA, USA) using the built-in rigorous column model. The Wilson equation was used for the evaluation of activity coefficients in the liquid phase and the Redlich-Kwong state equation for the vapour phase. The feed temperature was set at 30°C. Two prerequisites for the distillation design were 94 mass % of ethanol in the distillate and 99% recovery of ethanol from the feed.
RETROFIT
OF FERMENTATION
UNIT
Figure 2 shows the effect of the retrofit on the production capacity of the fermentation unit. The plant capacity can be doubled by adding the concentrated sugar feed in a mixing ratio of O-06. The maximum capacity in the range of mixing ratios investigated is almost eight times higher than in the reference case. Simultaneously, the output
0.2 0.3 Mixing ratio (-)
0.1
Fig. 2. Effect of retrofit on the production capacity of fermentation unit. Production capacity is given as a relative value of the production capacity of the reference case when the mixing ratio was equal to 0. The second curve represents the outlet ethanol concentration from the fermentation unit.
I6I
0.6
-
0
0.2
0.4
0.6
0.8
1
Split ratio C-1
sugar feed Fig. 3. Effect of distributing the concentrated on the volume of the cascade without cell recycle for a mixing ratio of 0.2. The split ratio represents the ratio of flow into the second stage to the total flow of concentrated feed. The fermenter volume is given in a relative form when it is compared to the volume at a split ratio equal to 0.
ethanol concentration from the fermentation unit increased from 1.3 to 7.5%. The results shown in this figure are not dependent on the fermentation unit set-up, but only on the stoichiometry and sugar conversion. In some previous studies dealing with multistage continuous ethanol fermentation, molasses was fed into several stages.‘9,20 In this case, the effect of distributing the concentrated sugar feed into the first and second stages on the retrofit design was studied. The last fermenter in the cascade could not be fed since this must operate at a low substrate concentration. It was found that feeding the second stage either did not affect the retrofit or was disadvantageous. Figure 3 illustrates an example where up to 60% of concentrated sugar feed could be fed to the second stage
Retrofit of ethanolfermentation
almost without any changes in the fermenter volume. This conclusion was only valid for mixing ratios smaller than 0.2. Above this value, the maximum fraction of the concentrated feed which could be fed into the second fermenter rapidly decreased. Consequently, multi-feeding was not considered in further calculations. As mentioned above, adding the concentrated sugar feed can increase the production capacity but this also means a higher flow rate. On the other hand, the higher amount of sugar fed into the system would require a longer residence time. Figure 4 illustrates that in some cases this demand might be satisfied without the installation of an additional fermenter. For example, only a 7% larger fermenter volume is required for doubling the production capacity in the system without recycle as follows from Figs 2 and 4. Besides the fermenter capacity, the centrifuge capacity must bc considered in the retrofit. The centrifuge capacity is determined by the suspension flow rate as well as by the suspension concentration. A higher concentration of solids implies a lower suspension flow rate that can be put through the centrifuge. Since in our case, adding the concentrated sugar feed means both higher flow rate and cell concentration, an increase in the centrifuge capacity is inevitable. The capacity of a centrifuge is generally a function of the solid size distribution and of the solid and liquid densities. In this study, the centrifuge capacity, CC, was then constant and represented a limit value of suspension flow rate for the yeast cell concentration close to 0. For the relation of the centrifuge capacity to the suspension flow rate, F,, we defined a quantity, called the centrifuge efficiency, r :
of low concentration
321
sugar solutions
According to vendor documentation on the commercial large scale centrifuges, with the capacity in the range of 75-200 m3/h of yeast cell suspension, the centrifuge efficiency is primarily dependent on the cell concentration but has very little dependence on the centrifuge size.*l The adaptation of these data gave the following polynomial relation of the centrifuge efficiency on the cell concentration (Fig. 5): ~=1-2~35x10-3c,-2-84x10-‘c2
X
(6)
Figure 5 shows that at a cell concentration of 100 kg/m3 the suspension flow rate amounts only to 50% of the centrifuge capacity. Figure 6 shows the relationship between the centrifuge capacity and the addition of concentrated feed assuming that the fermenter volume is the same as in the reference system. For a low mixing ratio, O-1, the relative increase in the centrifuge capacity is practically constant below a recycle ratio of 0%. This increase roughly corresponds to the flow of concentrated feed. On the other hand the required increase of centrifuge capacity is very high at a mixing ratio of O-4. As Fig. 6 shows, if the reference system has a recycle ratio above 0.4, which is relevant to continuous ethanol fermentations in industry, the system cannot be retrofitted by increasing the centrifugation capacity alone since this would require unrealistically large centrifugation capacities. The results above imply that a simultaneous increase in the fermenter volume and centrifuge capacity is required in some cases to retrofit the system, particularly when the concentrated sugar
2t 5
‘5 ;F
0.6
0.6
T
1
t_,,. 00
0.1
,
,
0.2
0.3
1 0.4
Mixing ratio (-)
Effect of the mixing ratio on the total volume of the Fig. 4. cascade without cell recycle. The fermenter volume is given as a relative value of the volume at a mixing ratio equal to 0 (reference case).
3 g z
0.4 0.2
u
0
Centrifuge Fig. 5. concentration.
efficiency
as a function
of yeast
cell
322
Milan Polakovi?,
Carl-Fredrik
feed ought to be added at a higher proportion. It was assumed that the total fermenter volume could be changed only by adding one or more fermenters of equal volume. The decision as to whether an extra fermenter should be used or not was based on a comparison of the costs of centrifugation and fermentation. We used an approximation that the costs of the installed fermenter to
centrifuge capacity were 3 : 1. This means that doubling the centrifuge capacity was economically equivalent to adding a fourth fermenter. Table 1 presents the retrofit results for all mixing ratios studied. It is shown here that a multiple increase in production capacity could be achieved with much less than a proportional increase in ferementer and centrifuge capacity. No additional fermenter installation is suggested at a mixing ratio of 0.1, which corresponds to a 150% increase in production capacity. The last column in Table 1 gives a multiplication factor characterizing the required increase in the centrifuge capacity of the retrofitted system. This factor varied between 37 and 156%. A certain increase in the centrifuge capacity might be managed without additional centrifuge installation depending on how far below the maximum capacity the centrifuge operates. Centrifuges, with the capacity relevant to the industrial ethanol fermentation, are manufactured in discrete sizes with a scale factor of 2~5.~~ This can give a significant difference between designed and utilized centrifuge capacity and can also give a space for the increased capacity required at retrofit. The economic evaluation of the retrofit of the fermentation unit followed from the results presented in Table 1. Since the retrofit was made with fermenters of equal volume, a linear approximation between the fermentation cost (excluding substrates) and total fermenter volume was appropriate. The centrifugation cost relative to the centrifuge capacity can also be expressed through a linear relationship.21’22 Thus, the comparisons of fermenter volumes and centrifuge capacities were also comparisons of the costs of the corresponding operations. Columns 3-6 of Table 2 compare the costs of fermentation and centrifugation with the refer-
r
0
02
0.4
0.8
RR,
0.8
1
t-1
Fig. 6. Relation between the centrifuge capacity and the mixing ratio in a cascade with recycle, expressed through the dependence of the relative centrifuge capacity (retrofitted versus reference system) on the cell recycle ratio in the referencc system with the mixing ratio as curve parameter: 1, mixing ratio 0.1; 2, mixing ratio 0.4. The dashed parts of the curves indicate the region where the recycle ratio in the retrofitted system is above the chosen, practical limit value of 0.9.
Table 1. Results of the retrofit of the fermentation unit (RR,, cell recycle ratio; c,;, cell concentration; MA, mixing ratio; V,,, fermenter cascade volume; CC,, centrifugation capacity). The subscripts o and R denote the reference and retrofitted cases, respectively RR, (CA
MR
0.75 (11.1)
0.1 0.2 0.3 0.4
V,,lv,
cxR
1 1.33 1567 2.33
32.6 43.3 56.2 61.0
CC&C”
RR,
082 0.72 0.83 082
Mandenius
1.37 1.70 2.22 2.56
Table 2. Cost related results of the retrofit of the fermentation unit (D,, mass tlow of ethanol obtained in the distillation unit). The subscript C denotes the conventional fermentation case. For the explanation of other symbols see Table 1 RR,
0.75
MR
v,
v,,
- vFu
D A
D,-0,
!!Ea
AV
D,
Do
cc, D _.A cc, D”
cc,
- cc,
D, -Do
VFR D A
cc,
v,
- v,,
cc,
- cc,
DR - Do
cc, D A
I.fx
EL
cc,
D*
DC
DC
DC
DC
0
4.95 2.52
1.08
o-12 0.16 0.24
1.92 1.81 1.63
1.02 1.19 l-14
0 0.1
0.37
0
0.51
0.22
1.23 0.46
;:; 0.4
0.3 O-27 0.3
0.13 0.10 0.20
0.39 0.37 0.33
0.2 1 0.24 0.23
0.37 0.34 0.37
D,-D, cc,
Retrofit of ethanol fermentation
ence case. All results are expressed per unit mass of 94% ethanol produced. While the odd columns show the relative costs for the retrofitted unit, the even columns represent the estimates of expenses per unit mass of extra ethanol obtained after the retrofit. The unit cost for both fermentation and centrifugation showed a significant decrease with the mixing ratio to 50% and less compared with the reference case. Both expenses decreased further with increasing mixing ratio. In general, the centrifugation cost varied over a smaller range than the fermentation cost. Even more convincing results were obtained after separating the cost of the reference system. In one case the fermentation expenses for additionally gained ethanol were equal to zero and their maximum value was 20% of the expenses in the reference system. The additional centrifugation cost varied only slightly and was somewhat more than 20% of that in the reference system. These figures demonstrate clearly the potential of the retrofit under study. Due to the decisive role of the substrate cost, it was of interest to compare the economics of the retrofitted plant with a conventional fermentation using substrates with high sugar content. This comparison is shown in the last four columns in Table 2. As a reference conventional fermentation was taken a system consisting of three fermenters in series fed with a substrate with a concentration of 164.3 kg/m3, and with a cell recycle of OS. The chosen substrate concentration corresponds to a total amount of sugar fed in the case of the mixing ratio of 0.4. The cell concentration at the outlet of the fermenter cascade was then 29.5 kg/m3. The advantages of the retrofit compared to the conventional fermentation were emphasized when the net expenses on additionally gained ethanol were evaluated. Although the extra centrifugation cost was about the same, the extra fer-
of low concentration
sugar solutions
323
mentation cost showed nearly the same positive outcome as in the previous comparison to the reference case (see columns 4 and 8 in Table 2).
RETROFIT
OF DISTILLATION
UNIT
Table 3 reviews the results of the distillation column calculations for different feed concentrations, XF, which correspond to the four different mixing ratios studied in the retrofit of the fermentation unit. As mentioned above, the amount of ethanol produced increased substantially with the mixing ratio, almost eight times compared to the reference case. As the distillate flow in none of these cases was above 10% of the feed flow value, the relative increase in the vapour flow at the column bottom was substantially smaller than the increase in the amount of distillate. The ratio of the bottom vapour flows was practically identical to the ratio of the reboiler duties. The relatively low consumption of heat in the reboiler at higher mixing ratios was also partially due to the decreased reflux ratio. The cooling duty of the condenser increased almost as sharply as the amount of distillate but was always smaller than the reboiler duty. The above results indicated clearly that adding the concentrated sugar feed would require the expansion of distillation capacities. The 30% increase in vapour flow at a mixing ratio of 0.1 could possibly be accommodated in the distillation column designed for the dilute ethanol feed but in any case the condenser capacity had to be increased significantly. The results presented in Table 3 were used for the cost evaluation of the distillation retrofit. Since the distillation column forms by far the largest portion of the total capital cost and the relative changes in the costs of the distillation column and reboiler were considered to be approximately
Table 3. Results of the retrofit of the distillation unit (MR, mixing ratio; &mass fraction of ethanol in the distillation feed; D/F, distillate to feed mass flow ratio; R, reflux ratio; D,, distillate flow; Q,,,,, Q,,,, reboiler and condenser duties, respectively; Vmi, bottom vapour flow). The subscripts o and R denote the reference and retrofitted cases, respectively
o-o
0.1 0.2 0.3 o-4
1.36 3.32 4.96 6.35 7.53
1.43 3.50 5.22 6.68 7.93
1 2.69 4.39 6.08 7.77
3.28 2-69 2.5 2.42 2.39
1 1.29 1.57 1.85 2.14
i.32 3.58 4.86 6.16
0.163 0.296 0.372 0.427 @469
324
Milan Polakovit, Carl-Fredrik Mandenius
Table 4. Cost-related
For the explanation MR
results of the retrofit of the distillation unit. The subscript C denotes the conventional of uther symbols see Table 3
e,,L
CL,< - dw,,
DtZD, L
v,,,
Do 0 0.1 0.2 o-3 0.4
K,,H - v,,,
D,-Do-D,-D,,
Do
Q,IIHV _I?1L!
Q,,v,,,,
Do
Do
D,
-
1.30
0.16 0.16 0.16
equal, a simplified but reasonably accurate cost evaluation was made. A linear approximation between the column cost including reboiler and vapour flow rate was considered. This approximation was compared with the empirical equation derived by Zacchi and Axelsson.7 In the range of the flow rates used, the deviation from the correct value was at a maximum at 10%. This deviation was compensated for by the condenser cost, which was otherwise neglected. The reboiler steam consumption was considered to be a key component of distillation operation cost. As this was proportional to the reboiler duty, a constant proportion between the distillation capital and operation costs was obtained in agreement with earlier results.’ Similarly as in the case of the fermentation unit, the economic consequences of retrofit were compared with the reference case as well as with conventional fermentation. The second and third columns of Table 4 arc rclatcd to the former process, while the last two columns of Table 4 to the latter. The second column of Table 4 shows that the estimated distillation cost per unit mass of final ethanol product already dropped to 48% at a mixing ratio of 0.1 and went further down to 28% at the mixing ratio of 0.4. The third column of Table 4 shows that the cost of retrofit per unit mass of additionally gained ethanol was independent of the mixing ratio and was only 16% of the cost in the reference plant. Obviously, the distillation cost per unit mass of ethanol (fifth column of Table 4) was higher than in the conventional fermentation. But, after the cost for the reference plant had been subtracted, the retrofit cost (last column of Table 4) showed a positive balance. The cost per unit mass of additionally gained ethanol showed a constant value of 60% of the cost in the conventional fermentation.
1.10
1
DISCUSSION
v,,
- v,,
i)h’(L D,
3.64 1.I4
0.16 0.36 0.30 0.28
- a,
case.
D,-D,,-D,-D,
n,-n,
e,L
ad
fermentation
D,
D,
0.60 0.60 0.60 0.61
AND CONCLUSION
The key factor in the evaluation of the economics of double sugar feed fermentation is the ratio of substrate costs. Two principal options can be defined here. First, that the costs of both substrates on the total mass basis are comparative and secondly, that the low content sugar substrate is substantially cheaper. Both alternatives are reasonable. The cost of cellulose hydrolysate, even from waste materials, is rather high due to the high cost of conversion of cellulose to fermentable sugars.” On the other hand, some wastes containing fermentable sugars like spent sulphite liquor or whey, arc rather cheap, especially when they are further utilized on the production site. For example, the calculated price of spent sulphite liquor in the pulp plant is 0.8 USS per m3 of liquor. This low price of the liquor makes the production of technical ethanol profitable in spite of the low sugar content. The price of molasses is currently about 325 US$ per tonne in Sweden and the transportation cost amounts to 75 US$ per tonne resulting in a total cost of approximately 200 US$ per tonne of molasses. The economic effects of the retrofit presented in Tables 2 and 4 are related to the above two substrate cost ratios in a different way. For cellulose hydrolysate, the bcncfits of the retrofit are better demonstrated if the costs are compared to the reference process. The retrofit improves the economy of the whole process and brings down the cost of ethanol produced in this way. For a cheap substrate, it is more relevant to compare the costs with conventional fermentation since, after adding a concentrated sugar feed, the total cost of produced ethanol will increase but this will be economically justified while the ethanol cost remains lower than the cost achieved in a conventional fermentation plant.
Retrofit of ethanol fermentation of low concentration
A relatively small increase in the mixing ratio of sugar feeds resulted in a substantial increase in the plant production capacity which could be accommodated in the existing plant facilities with minor modifications. A plant designed for a certain ratio of sugar feeds would thus allow rather flexible operation to accommodate seasonal fluctuations in feedstock amounts or price fluctuations.
ACKNOWLEDGEMENTS This work was supported by a grant from the Swedish Ethanol Foundation and the National Board for Technical Development (NUTEK). Professor Guido Zacchi is kindly acknowledged for his advice on ethanol distillation. We are grateful to Mr Hans Axelsson, from Alfa Lava1 Separation AB for the centrifuge documentation and personal information. We thank MS Eva Larsson, from MoDo Paper AB, t)rnskiildsvik and MS Tngrid Ascard, from Sockerbolaget AB, Malmii for providing information on the costs of spent sulphite liquor and molasses, respectively.
7.
8. 9.
10.
1 I.
12.
13. 14.
15.
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