- Email: [email protected]
Energy considerations for a SSF-based softwood ethanol plant
Available online at www.sciencedirect.com
Bioresource Technology 99 (2008) 2121–2131
Energy considerations for a SSF-based softwood ethanol plant Anders Wingren *, Mats Galbe, Guido Zacchi Department of Chemical Engineering, Lund University, P.O. Box 124, SE-221 00 Lund, Sweden Received 8 September 2006; received in revised form 23 May 2007; accepted 24 May 2007 Available online 27 September 2007
Abstract Ethanol can be produced from softwood by steam pretreatment followed by simultaneous saccharification and fermentation (SSF). However, the final ethanol concentration in the SSF step is usually rather low (around 4 wt%) and as a result the energy demand in the downstream processing will be high. In an effort to reduce the energy consumption various alternatives for the downstream processing part of the process were evaluated from a technical-economic standpoint. With experimental data as a basis, the whole process was modelled using the commercial flowsheeting program Aspen Plus. The results were used in the subsequent economic evaluation, which was performed using Icarus process evaluator. A base case configuration, consisting of three thermally coupled distillation columns and multiple-effect evaporation was established. For a feed containing 3.5% ethanol (w/w) to the distillation step, the overall process demand for steam was estimated to be 19.0 MJ/L ethanol and the ethanol production cost 4.14 SEK/L (0.591 USD/L). Different alternatives were considered, such as integration of a stripper with the evaporation step, increasing the number of evaporation effects and the application of mechanical vapour recompression to the evaporation step. Replacement of evaporation with anaerobic digestion was also considered. Among these alternatives, evaporation using mechanical vapour recompression and the anaerobic digester alternative both resulted in significantly lower energy demand than the base case, 10.2 and 9.8 MJ/L, respectively, and productions costs of 3.82 (0.546 USD/L) and 3.84 SEK/L (0.549 USD/L). Ó 2007 Elsevier Ltd. All rights reserved. Keywords: Ethanol; Economics; Process; SSF; Energy optimization
1. Introduction We are now having to face the fact that the world’s resources of oil will eventually be depleted and that alternative fuels will have to be found. The optimal replacement fuel should be cheap and it should be possible to produce it in large quantities from a renewable source. Furthermore, the life cycle of the fuel should be environmental friendly. It would be an advantage if this fuel could also be used directly in existing motor vehicles and the same distribution system could be used as for petrol. Bioethanol exhibits almost all these desirable properties. It can be produced from lignocellulosic feedstocks, which are cheap, renewable and available in large quantities. If bioethanol *
Corresponding author. Tel.: +46 46 2228297; fax: +46 46 2224526. E-mail addresses: [email protected] (A. Wingren), Guido. [email protected] (G. Zacchi). 0960-8524/$ - see front matter Ó 2007 Elsevier Ltd. All rights reserved. doi:10.1016/j.biortech.2007.05.058
were to replace petrol the emission of greenhouse gases would be reduced by more than 85% when considering the whole fuel cycle (Bergeron, 1996). Ethanol also has the advantage of being easy to distribute. Gasoline–ethanol blends containing up to around 20% ethanol can readily be used in modern spark-ignited combustion engines, while flexible fuel vehicles can run on conventional petrol, ethanol or combinations of the two. Various studies have concluded that it is possible to produce ethanol from lignocellulosic raw materials in an economically feasible way (Aden et al., 2002; Lynd et al., 1996; von Sivers and Zacchi, 1995; Wingren et al., 2004). In Sweden softwood is considered to be the raw material with the highest potential for ethanol production. It is abundant and has a high content of carbohydrates that can be hydrolysed to fermentable sugars. In previous techno-economic evaluations of softwood as raw material, a process consisting of steam pretreatment followed by
2122
A. Wingren et al. / Bioresource Technology 99 (2008) 2121–2131
enzymatic hydrolysis and fermentation was studied (von Sivers and Zacchi, 1995; Wingren et al., 2004). In this process the lignin-rich residue is a valuable co-product that can be dried, pelletized and sold as a solid fuel. However, a major part of the solid residue is needed for the production of process steam. Compared with sugar-based or starchbased processes the feed to the distillation step in a future lignocellulosic plant is likely to be more dilute, and as a consequence, the downstream processing will be energy demanding and thus costly. To increase the income from this co-product the energy demand in the process has to be reduced. Several options are available. (i) A higher substrate load in the saccharification and fermentation steps increases the ethanol concentration in the feed to the downstream processing steps. Thus the energy demand in the distillation and evaporation steps decreases. (ii) The fresh water added in the saccharification and fermentation steps can be replaced with recirculated process streams such as the stream before distillation or the stillage stream after distillation (Alkasrawi et al., 2002). The latter is an option widely employed in sugar-based and starch-based ethanol processes (Shojaosadati, 1996; St. Julian et al., 1990). (iii) The energy demand can be reduced by energy integration where the use of secondary steam and waste heat is optimized. (iv) Another option is to optimize the energydemanding process steps directly, e.g. by employing more advanced distillation or evaporation systems or by replacing these steps with less energy-intensive unit operations. The main objectives of downstream processing are to produce fuel ethanol, to recover valuable co-products and to produce wastewater that can either be recycled in the process or easily be disposed of. Although there are several options, distillation is almost the only unit operation used for concentration of dilute ethanol streams in commercial ethanol processes. Several treatment methods are available for the stillage and an excellent review paper on this topic has been written by Wilkie et al. (2000). In sugar- and starch-based ethanol plants evaporation, anaerobic digestion, membrane filtration and direct application to land have been used as methods of stillage treatment. Combinations of these methods are also common. Regarding lignocellulosic feedstocks, only limited data regarding the treatment of the stillage have been published. Evaporation is an energy-demanding unit operation but it produces a syrup that can be combusted, and clean condensates that can be recycled or treated further before being disposed of. Alkasrawi (2004) studied the effect of recycling evaporation condensates in a simultaneous saccharification and fermentation (SSF) process, and Larsson et al. (1997) performed similar studies on a separate hydrolysis and fermentation (SHF) process. Their conclusions were that a major part of the fresh water added to the process could be replaced with condensates without affecting ethanol yield or production rate. Anaerobic digestion, followed by an aerobic step, is an interesting option since it has a low energy demand and produces methane, which can be used for steam production. Callander et al. (1986) studied
anaerobic digestion of softwood stillage with a COD and BOD content of 25 000 and 13 000 mg/L, respectively. A removal of 86% of the COD and 93% of the BOD was achieved. The effect of recycling this purified stream was not investigated. This study deals with the ‘‘ethanol-from-wood’’ process from an energy perspective. A base case is established and used for comparison with alternative process configurations in which energy-saving options, such as multiple-effect evaporation and integration of evaporation and distillation are employed. The use of mechanical vapour recompression and replacement of evaporators with an anaerobic digester have also been investigated. Since downstream processing is highly energy demanding, distillation and evaporation were studied in detail, but as the energy requirement of a plant of this kind cannot be optimized by the investigation of a single step, the fully integrated process was evaluated. 2. Process description 2.1. Pretreatment and SSF Fig. 1 shows the flowsheet for the process on which the simulations were based on. Fresh wood chips of spruce are impregnated with sulphur dioxide (2 wt% of the liquid in the raw material) and thereafter steam pretreated at 215 °C for 5 min. The slurry is flash-cooled by pressure reduction to 4 bar and then to 1 bar, after which it is fed to the SSF step. The pH-value is adjusted with lime to about 5 and the slurry diluted to a concentration of water-insoluble solids (WIS) of 8.4%. After temperature adjustment to 37 °C and addition of enzymes, corresponding to 15 filter paper units per gram WIS, and nutrients, the conversion starts and is completed in 72 h. The process is assumed to be operated batchwise with a total cycle time of 84 h. Process data for these steps are based on experimental work carried out in a process development unit (PDU) at the Department of Chemical Engineering, Lund University, Sweden. The dry raw material consists of 45.0% glucan, 12.6% mannan, 2.6% galactan, 7.1% pentosans, 28.1% lignin and the remainder is made up of acetyl groups, extractives and ash. The water content is 50%. In the pretreatment step 15% of the glucan is converted to glucose and 10% to byproducts such as 5-hydroxymethyl furfural. Of the mannan, galactan, arabinan and xylan, 70% is converted to monomeric sugars and 25% to by-products. Around 5% of the lignin is solubilized. The mass recovery, based on dry weight, in this step is 90% of the dry raw material. In the subsequent SSF step 94% of the glucan is hydrolysed. Of the glucose and mannose, 92% is converted to ethanol and the remaining fraction is converted to by-products or remains unfermented. All other sugars fed to the SSF step are unconverted. The major difference between the experimental work and the simulation is the mode of operation in the SSF step which was fed-batch in the experimental work
A. Wingren et al. / Bioresource Technology 99 (2008) 2121–2131
Wood
Feedstock handling
Pretreatment
SSF
2123
Ethanol
Distillation Stillage
Evaporation
Solid-liquid separation
Pelletizing Drying
Solid fuel
Solids
Incineration
Steam
Syrup
Other
Fig. 1. Flowsheet for the ethanol-from-wood process.
but assumed to be batch in the simulation and in the economic evaluation. Also, baker’s yeast (Saccharomyces cerevisiae) was added to a concentration of 5 g/L in the experiments. In a full-scale process it is more likely that the yeast would be produced in-house using some of the sugar available in the hydrolysate as substrate. In the simulation it was therefore assumed that the yeast is produced directly in the fermentors to a concentration of 2 g/L prior to the start of SSF. 2.2. Distillation and evaporation The flowsheet for this step is shown in Fig. 2. The base case consists of two strippers and a rectifier with top stage pressures of 3, 1.25 and 0.35 bar. The stream from the SSF step is preheated and then divided between the two strippers. Overhead vapour from the stripper at 3 bar is used as the heat source in the reboiler of the second stripper before being fed to the rectification column. Overhead
vapour from the second stripper, as well as some primary steam, is used as the heat source in the reboiler of the rectification column. The feed to the two strippers consists of the unfiltered mash containing not only water and ethanol, but also solids such as lignin and yeast. Therefore, it can be expected that the tray efficiency will be quite modest, and in this study a Murphree efficiency of 50% was used. The rectifier, operating without solids, is expected to have a higher efficiency, 75%. The recovery of ethanol was assumed to be 99.5% in all three columns and the distillate from the rectifier was assumed to consist of ethanol at 94% w/w. For the specifications given above for the rectifier the reflux ratio is dependent on the number of trays. A higher reflux ratio reduces the number of trays, makes the column diameter larger and requires a higher reboiler duty. Thus, the optimal reflux ratio, i.e. that resulting in the lowest overall cost, is a trade-off between capital cost and energy cost. Simulations were carried out to study how the number of trays
Ethanol (94% w/w)
Stripping
Evaporation
Rectification
Centrifugation
Syrup (50% DM)
From SSF
Fig. 2. Flowsheet for the distillation and the evaporation steps in the base case. Units using primary steam or secondary steam from other process steps, i.e. the dryer or the pretreatment are shaded. Dotted lines indicate vapour phase.
2124
A. Wingren et al. / Bioresource Technology 99 (2008) 2121–2131
affected the reflux ratio and the reboiler duty, and it was concluded that a reasonable number of trays was 35 in the rectifier with a reflux ratio of 2.4. The two strippers are equipped with 25 trays each. The stillage is sent to a liquid–solid separation step consisting of two decanter centrifuges connected in parallel. The concentration of WIS in the solid fraction is assumed to be 30%. In the evaporation step the liquid is concentrated to 50% dry matter (DM) and part of the condensate can be recycled in the SSF step. A multiple-effect evaporator system consisting of five effects with a forward feed arrangement was used in the base case configuration. The pressure in the first effect is 3 bar and in the last 0.2 bar. The pressures in effects 2–4 were adjusted to achieve roughly the same difference in temperature between the hot and cold side of all effects. The boiling point elevation for wood hydrolysates has been determined experimentally (Larsson et al., 1997) and was accounted for. Overall heat transfer coefficients were varied from 2000 to 600 W/m2 °C depending on the temperature and concentration of the slurry being evaporated. Outgoing condensate is used for preheating of the feed to the strippers. 2.3. Steam generation and solid fuel production The concentrated stream from the evaporator and the solid stream from the centrifuges are mixed and fed to the drying step where the material is dried to 85% DM. The equipment is assumed to be a steam dryer in which superheated steam is used as the drying medium. Primary steam is required at two pressure levels, 20 and 12 bar, and the dryer works at 4 bar, which means that secondary steam is available at a high temperature and, thus, the net energy demand per unit evaporated liquid is low. Part of the dried solid residue is sent to the boiler where fresh
steam for the process is generated. Excess solid residue is pelletized and sold as a solid fuel. 2.4. Alternative process configurations Different process configurations for the distillation and evaporation steps were investigated. The process steps on the upstream side of these steps are unaffected. The following options were considered. 2.4.1. Increasing the number of effects in the evaporator The number of effects in the evaporation step was increased from the base case value of five up to eight effects. The pressures in the first and the last effect were the same as in the base case. The highest possible number of effects is determined by the overall temperature difference over the evaporation step. 2.4.2. Integration of the stripper with evaporation A single stripper working at 1 bar is integrated in the evaporation train with five effects, see Fig. 3. The steam from one of the evaporation effects is used as the heating medium in the reboiler of the stripper. The vapour overhead from the stripper is used as the heating medium in the following effect in the evaporation step. Thus, in this case the rectifier, which is operated at atmospheric pressure, must use primary steam as its heating source. 2.4.3. Mechanical vapour recompression applied to the evaporator A large portion of the energy supplied to a traditional multiple-effect evaporator ends up as latent heat in the vapour phase leaving the last effect in the evaporator. This vapour is normally condensed using cooling water. Another option is to compress the vapour, thereby raising the
Ethanol (94% w/w)
Stripping Rectification To drying
Evaporation
From SSF
Fig. 3. Flowsheet for the stripper integrated with the evaporation unit.
A. Wingren et al. / Bioresource Technology 99 (2008) 2121–2131
temperature to a level at which the latent heat will be available. The vapour can then be used as a heating medium to replace most of the primary steam. When compression is carried out by a mechanical compressor the process is referred to as mechanical vapour recompression (MVR). An electrical motor or a steam turbine provides the thrust to the compressor. Another option is to thermally recompress the vapour by use of a steam ejector. This process is called thermal vapour recompression. This study was limited to the use of MVR with an electrical motor. The economy is dependent on the power demand of the compressor and this is governed by the vol-
2125
umetric flow rate of the vapour and the pressure increase over the compressor. The pressure increase governs the driving force in the evaporator (i.e. the temperature difference), which affects the area needed for heat transfer. Thus, as the electricity consumption increases, the evaporator area decreases. The purpose of this study was not to fully optimize this system, but to investigate the potential of MVR in an ethanol plant. Compression from 3 to 4 bar in a single step (two effects in parallel) was assumed providing a temperature difference of 10 °C, see Fig. 4. In the MVR unit the slurry is concentrated to 20% and thereafter to 50% in a small 4-effect evaporator. The reason for not
Ethanol (94% w/w)
Stripping Compresser
Rectification
To drying
Evaporation
From SSF
Fig. 4. Mechanical vapour recompression applied to the evaporation step.
Ethanol (94% w/w)
Rectification
Stripping
To drying
From SSF
Biogas to boiler
Anaerobic digestion
Aerobic step
Purified water
Fig. 5. Evaporation replaced by anaerobic digestion followed by aerobic digestion.
2126
A. Wingren et al. / Bioresource Technology 99 (2008) 2121–2131
concentrating the slurry to 50% in the MVR unit directly is the boiling point elevation, which is 7 °C at 50% but only 1 °C at 20% for this kind of liquid (Larsson et al., 1997). 2.4.4. Anaerobic digestion The organic material in the stillage is converted through anaerobic digestion (AD) to biogas consisting of methane and carbon dioxide, see Fig. 5. The performance of such a system is dependent on a number of parameters such as the composition of the feed, residence time, temperature, etc. Since very limited data regarding the performance of this kind of system have been published only a rough estimate could be made based on data from PURAC AB, Lund, Sweden. The liquid part of the stillage is cooled to the operating temperature of the anaerobic digester (55 °C) by heat exchange with the feed to the distillation step. The production of methane is assumed to be 0.35 m3/kg COD consumed and in the anaerobic step around 50% of the COD is removed. In the subsequent aerobic step most of the remaining COD is removed and this water can then be either recycled to the SSF step or disposed of. The residence time in the anaerobic step is 20 days and in the aerobic step 4 days. Nitrogen and phosphor are added at a rate corresponding to 26 kg ammonia/h and 23 kg phosphoric acid/h. The biogas produced is sent to the steam boiler. Sludge will also be produced, especially in the aerobic step. However, the amount and composition are associated with great uncertainties. Sludge can be dewatered and thereafter landfilled or incinerated. These aspects were not studied in detail in this study, and it was assumed that the contribution of sludge to the overall production cost is minor and was thus neglected. The process configurations studied were all based on the same fundamental process design. Primary steam was assumed to be available at 25, 12 and 6.9 bar and secondary steam was used to replace fresh steam when possible. The plant considered is a grass-root plant, designed to process 200 000 tonnes of raw material annually on a DM basis. Annual on-stream time was assumed to be 8000 h. 3. Methods The process simulations were carried out using Aspen Plus version 11.1 from Aspen Technology. All process steps described above were simulated using built-in modules for reactors, heat exchangers, pumps, etc. Distillation columns were simulated rigorously, i.e. using the RADFRAC module. Results from the process simulations were used in the sizing of the process equipment. Capital costs for standard equipment such as heat exchangers, columns, evaporators, vessels and pumps were estimated using Icarus process evaluator (IPE), version 12.0 from Aspen Technology. Auxiliary equipment such as piping, electrical equipment and instrumentation, as well as indirect costs arising from engineering, freight, construction indirects and contingency, were also estimated using IPE. The equipment cost estimated by IPE was compared with quo-
tation costs obtained from Swedish suppliers of the same kind of equipment and showed good agreement. The pretreatment reactor was assumed to work like a StakeTech reactor and a quotation provided from Stake Technology was used in the cost estimation. Cost data for the pellet machine, anaerobic digester, compressor for the mechanical vapour recompression and the dryer were also obtained from vendors, and data for the steam boiler were obtained from the literature. Working capital was accounted for as the cost of raw material and chemicals for 10 days production, the value of final products for 30 days production, 30 days of operating expenses and accounts receivable to allow a 30-day payment credit for customers. The annual capital cost was determined by multiplying the fixed capital investment by an annuity factor of 0.103 corresponding to a 15-year pay-off time and an interest rate of 6%. This should be regarded as a favourable bank loan. A higher interest rate and/or a shorter pay-off time would result in a higher annuity factor and thus a higher annual capital cost. The cost of the raw material was assumed to be 90 SEK/MW h, while the income from the solid fuel coproduct was 140 SEK/MW h. Costs used in this evaluation are summarized in Table 1. The purpose of the process simulations and the economic evaluation was to compare different process configurations on an equal basis and not to obtain an absolute ethanol production cost. Table 1 Costs used in the economic evaluation Type
Cost
Unit
Raw material Wooda
90
SEK/MW h
Chemicals SO2 CaO Defoamer NaOH (50%) (NH4)2HPO4 MgSO4 Æ 7H2O NH3 (29%) H3PO4 (50%) Enzymes
1.5 0.7 20 2.0 1.5 4.4 2.6 5.0 19
SEK/kg SEK/kg SEK/kg SEK/kg SEK/kg SEK/kg SEK/kg SEK/kg SEK/106 FPU
By-product income Solid fuelb Carbon dioxide
140 0.03
SEK/MW h SEK/kg
Utilities Electricity Cooling water Process water
250 0.14 1.40
SEK/MW h SEK/m3 SEK/m3
Other costs Labour
500 000
Insurance Maintenance
1 2
SEK/employee year (25 employees assumed) % of fixed capital per year % of fixed capital per year
Capital costs Annuity factor
0.103
a b
0.42 SEK/kg DM. 0.79 SEK/kg DM.
A. Wingren et al. / Bioresource Technology 99 (2008) 2121–2131
2127
steam pretreatment of hardwood using roughly the same process conditions as in this study. They concluded that approximately 50% of the sulphur added in the impregnation step remained in the slurry following steam pretreament. Only a small fraction (7% of the added sulphur) was bound to the water-insoluble fraction, and the sulphur in the liquid phase was mainly in a non-volatile form such as lignosulphonates. This result clearly indicates that most of the sulphur added in the pretreatment step in the evaluated plant will be found in the syrup stream leaving the evaporation unit and not in the fibrous stream. Since the excess solid residue is intended to be sold as a solid fuel, a low sulphur content, a high heating value and a low ash content are desirable. The syrup can be combusted directly instead of mixing it with the fibrous stream prior to drying. Part of the dried fibrous stream can also be incinerated as well if necessary. Further research is needed in this area. Table 2 lists the energy demand in the different process steps for the various alternatives. In the base case configuration the process steps with the highest net energy demand are distillation and evaporation. These two steps require 29.6 MW steam. Pretreatment and drying require high pressure steam but also produce significant amounts of secondary steam (16.3 MW), which can replace some of the primary steam needed in the distillation and evaporation steps. The total energy demand in the form of primary (boiler-generated) steam is thus 38.8 MW or 19.0 MJ/L ethanol. This steam is generated by incinerating 61% of the 12 000 kg/DM h of solids from the drying step thus leaving 4650 kg DM/h for sale. Fig. 6 shows an energy flow diagram of the base case configuration. The energy bound in the raw material is divided mainly between ethanol and solid fuel. Of the 143 MW originating from the raw material, 47.9 MW or 34% ends up as ethanol. The excess solid fuel represents 28.9 MW or 20%. Thus,
4. Results and discussion With experimental data as the basis the process was simulated in Aspen Plus to determine unknown process parameters such as temperatures, size and composition of process streams, as well as utility requirements in the different process steps. For a process with 8.4% WIS in the SSF step the concentration of ethanol in the feed to the distillation step is 3.5% (w/w) and the concentration of DM in the stream being fed to the evaporation or anaerobic digestion step is 3.3%. Ethanol is produced at a rate of 5813 kg/h, which corresponds to an overall ethanol yield of 71% of the theoretical based on the hexoses fermentable by baker’s yeast (i.e. glucan and mannan) in the raw material. The total amount of solid residue before the steam generation step, 12 000 kg DM/h, is the same for all process configurations except the AD case where the stillage stream, having a flow of dry matter of around 4200 kg/h, is fed to the anaerobic digester. The flow of solid residue before steam generation is thus lower, 7800 kg DM/h. The methane production in the AD case was estimated to be 667 kg/h. There is a significant difference between the solid residue in the AD case and in the other process configurations. In the AD configuration, the residue consists of around 80% lignin and the remaining fraction is made up of nondegraded polysaccharides, yeast, etc. In the other cases the syrup from the evaporator has been mixed with the fibrous stream prior to drying. Since the syrup consists of non-volatile solubles such as glycerol, dissolved lignin, sugars, ash, inorganic compounds, etc., the lignin content of the solid residues decreases and is estimated to be only 55%. This solid stream will have a lower heating value than the AD residue and possibly also have a lower quality as a fuel due to its higher ash content. The syrup may also contain more sulphur than the fibrous stream. (Gregg and Saddler (1996) calculated a sulphur balance for SO2-catalysed
Table 2 Overall heat duties (MW) for various process configurations, as well as the contribution from the individual process steps to the overall heat duty BC
Steam required (MW) Pretreatment Distillation Evaporation Drying
MEE8
Intstripp
MVR
HP
LP
HP
LP
HP
LP
HP
10.0
4.6 15.8b 13.8
10.0
4.6 15.9 7.7
10.0
4.6 9.7 18.4
10.0
10.9
Steam generated (MW) Pretreatment Drying
10.8 6.2 10.1
Net energy demanda
20.9
Thermal energy boiler (MW)
38.8
17.9
11.3 6.2 10.1
20.9 32.9
12.0
37.3
LP 4.6 8.4 2.5
10.9 6.2 10.5
21.3
AD
16.0
HP
LP
10.0
4.6 11.3
8.6 6.2 10.1
20.9 20.9
0.9
6.2 8.4 18.6
1.4
20.0
Data are presented for 8.4% WIS in the SSF step. BC = base case, MEE8 = multiple-effect evaporation with eight effects, Intstripp = stripper integrated with evaporator, MVR = mechanical vapour recompression, AD = anaerobic digestion. HP: High pressure steam 25 and 12 bar, LP: low pressure steam 4 and 1 bar. a Difference between steam required and steam generated. b Of which 7.4 MW is needed for preheating of the feed.
2128
A. Wingren et al. / Bioresource Technology 99 (2008) 2121–2131 Solids 55.9 (54.3)
Losses and cooling 3.6
Pretreatment
Hydrolysate 129 (126)
SSF
Broth 129 (126)
Flash condensate18.3 (17.3)
Wood 143 (143)
Excess solid fuel 29.3 (28.9) Drying
37.3 (22.1)
Fuel 46.4 (45.8)
Stillage 93.9 (76.5)
Ethanol 47.9 (47.9)
Solid sep.
Distillation
Condensates 5.2 (1.2)
Preheating 6.3
Condenser 18.4
Evaporation
Solid fuel 20.9 (20.7)
Condenser 5.7
Incineration
Losses 7.7 (1.8)
LP steam 15.8
HP steam 10.0 LP steam 4.6 LP steam 6.2
Steam distribution
LP steam 13.8 HP steam 10.9 LP steam 10.1 Steam 38.8
Fig. 6. Energy flows (in MW) in the base case configuration at 8.4% WIS in the SSF step. Raw material intake is 25 000 kg DM/h. The values given are energy flows (based on heat of combustion) calculated relative to a reference state of 25 °C and water in liquid phase. Values in brackets are without water. Heat flows, representing the net flow of energy, are shown as dotted lines. Balances may be inconsistent due to rounding-off errors and due to the exclusion of minor streams.
the overall energy efficiency is 54%. However, the process also requires 5.5 MW electricity. If this figure is included, assuming an electricity-to-fuel ratio of 0.3, the energy efficiency becomes 48%. Thirty-two percent of the energy in the raw material is used for steam production, while 12% ends up in the flash steam from the pretreatment step. The latter figure is associated with large uncertainties due to the difficulty in determining the composition of the flash steam. In the PDU, around 90% of the dry material is recovered as dry matter in the slurry following steam pretreatment. The fraction of the material not recovered as dry matter is likely to be volatile substances formed during pretreatment as with extractives from the wood. The exact composition is, however, unknown. All the alternative process configurations result in a lower energy demand than the base case. The energy demand is reduced as the number of effects in the evaporation step is increased. When eight effects are used the energy demand is reduced to 32.9 MW, while the case with an integrated stripper results in a minor reduction in energy demand. The two remaining alternatives, the MVR case and the AD case, both reduce the overall energy demand significantly. The use of MVR results in an overall energy demand of 20.9 MW and in the anaerobic digester case 20.0 MW. The electricity needed for the compressor in the former alternative is 2.3 MW. The energy demand in the distillation step is lower for the MVR and AD cases than for the base case despite the fact that the same process configuration is used in the distillation step. This is explained by the fact that these two steps are integrated, and if the process conditions are changed in one of the steps this will affect the other. In the case employing MVR the condensate stream, having a temperature of
144 °C, can provide all the energy needed for preheating of the feed to the strippers. When an anaerobic digester replaces the evaporation step, the feed to the strippers is preheated by the stillage, which can also supply essentially all the energy needed for preheating. In fact, the MVR and anaerobic digestion cases are so efficient that all the secondary low-pressure steam generated in the dryer and the pretreatment step cannot be utilized in the distillation and the evaporation steps. However, 4.6 MW can be used to preheat the wood in the pretreatment step to 120 °C. Despite this, the mechanical vapour recompression case has a small surplus of secondary steam corresponding to 0.9 MW, which is condensed by cooling water. The evaluated process configurations are summarized in Table 3. An overall capital investment of 889 million SEK was estimated for the base case process and the ethanol production cost was estimated to be 4.14 SEK/L (100% ethanol). The cost of the raw material corresponds to 1.41 SEK/L and the income from the solid fuel co-product is 0.50 SEK/L. The ethanol production cost decreases as the number of effects in the evaporation step is increased. However, the difference in cost between seven effects and eight effects is minor (the latter being 4.06 SEK/L), see Table 3. Increasing the number of effects further would thus not result in a lower ethanol cost but would result in a very low driving force (temperature difference) in the individual effects. Integration of the stripper in the evaporation step does not result in any major cost reductions. The drawback with this configuration is the temperature drop over the stripper, which reduces the available temperature difference over the evaporation step. The result is increased heat transfer area in the evaporators compared with the base case.
A. Wingren et al. / Bioresource Technology 99 (2008) 2121–2131
2129
Table 3 Summary of energy demand and costs for the process configurations evaluated with 8.4% WIS in the SSF step
Heat requirement (MJ/L ethanol) Total capital investment (MSEK)
BC
MEE6
MEE7
MEE8
Intstripp
MVR
AD
19.0 889
17.7 896
16.8 903
16.1 912
18.3 895
10.2 890
9.8 853
Costs (SEK/L ethanol) Raw material Solid fuel Capital Electricity Chemicals Enzymes Other
1.41 0.50 1.51 0.19 0.26 0.60 0.67
1.41 0.55 1.52 0.18 0.26 0.60 0.67
1.41 0.59 1.53 0.18 0.26 0.60 0.68
1.41 0.62 1.55 0.18 0.26 0.60 0.68
1.41 0.53 1.52 0.19 0.26 0.60 0.68
1.41 0.86 1.51 0.26 0.26 0.60 0.65
1.41 0.70 1.45 0.18 0.28 0.60 0.63
Total production cost (SEK/L)
4.14
4.10
4.07
4.06
4.12
3.82
3.84
Total production cost (USD/L)
0.591
0.586
0.581
0.580
0.589
0.546
0.549
BC = base case, MEE8 = multiple-effect evaporation with eight effects, Intstripp = stripper integrated with evaporator, MVR = mechanical vapour recompression, AD = anaerobic digestion.
4.5
Ethanol production cost (SEK/L)
The MVR case results in an overall ethanol production cost of 3.82 SEK/L. The income from the solid fuel coproduct is 0.86 SEK/L. Although the cost of the evaporation system increases, the total capital investment increases only slightly due to the cost reductions in the steam generation step. The cost of the electricity is higher in this case due to the power requirement of the compressor. Almost as low production cost, 3.84 SEK/L, is achieved with the process alternative using AD. This results in a lower capital cost as well as a higher income from the solid fuel co-product than the base case. An analysis was carried out to investigate how sensitive the ethanol production cost is to variations in the selling price of the solid fuel co-product. The price was varied from 50 SEK/MW h to 250 SEK/MW h and the results are shown in Fig. 7. The largest cost reduction is seen for the MVR case since this configuration results in the highest production of solid fuel, and is thus more sensitive to price changes. For the same reason, the base case configuration has the lowest sensitivity. At 200 SEK/MW h the cost difference between the case with the highest production cost (base case) and the case with the lowest production cost (MVR), is 0.74 SEK/L ethanol. However, if the selling price is 50 SEK/MW h this difference is only 0.17 SEK/L. When the process configurations are evaluated assuming a higher concentration of WIS in the SSF step, the production cost is reduced for all alternatives, see Fig. 8. An increased concentration of WIS results in increased ethanol concentration in the feed to the distillation step. At 10% and 12% WIS the ethanol concentration is 4.2% and 5.1%, respectively. The concentration of DM in the feed to the evaporator and the anaerobic digester also increases, being 4.0% and 4.9% for 10% and 12% WIS in SSF, respectively. The base case and the case with eight effects show a larger decrease in ethanol production cost than the AD and MVR configurations. This can be explained by the higher energy demand in the multiple-effect evaporation cases. The energy demand in downstream processing decreases when the concentration of WIS is increased in the SSF step,
BC MEE8 MVR AD
4.0
3.5
3.0 0
50
100
150
200
250
300
Price of solid fuel (SEK/MWh)
Fig. 7. Ethanol cost as a function of the price of the solid fuel co-product. Data for 8.4% WIS in the SSF step. BC = base case, MEE8 = multipleeffect evaporation with eight effects, MVR = mechanical vapour recompression, AD = anaerobic digestion.
and this effect is more noticeable for process configurations having a high energy demand than for energy-optimized systems. The ethanol production cost at 12% WIS is 3.73 SEK/L for the base case configuration and the lowest cost at this WIS concentration was obtained for the MVR configuration, being 3.57 SEK/L. As stated above, the evaluation is based on data obtained from experiments carried out at 8.4% WIS (fedbatch) in SSF, and the evaluation of the 10% and 12% WIS cases is based on the assumption that the yield, as well as the productivity, will be the same as in the 8.4% case. S. cerevisiae is a robust yeast that can tolerate ethanol concentrations well above 5% (corresponding to 12% WIS in SSF). However, wood hydrolysates contain other compounds which may be toxic to the yeast, as well as to the enzymes, and thus the ethanol yield as well as the
2130
A. Wingren et al. / Bioresource Technology 99 (2008) 2121–2131 4.8
4.2 BC MEE8 AD MVR
4.6
Ethanol production cost (SEK/L)
Ethanol production cost (SEK/L)
4.1
4.0
3.9
3.8
3.7
3.6
3.5 8
8.4% WIS 10% WIS 12% WIS
4.4
4.2
4.0
3.8
3.6
9
10
11
12
13
% WIS in SSF
Fig. 8. Ethanol production cost as a function of WIS for the process configurations evaluated. BC = base case, MEE8 = multiple-effect evaporation with eight effects, MVR = mechanical vapour recompression, AD = anaerobic digestion.
productivity may be reduced as the concentration of WIS increases. It is the authors’ belief, however, that an optimized fed-batch process at 12% WIS with a yeast adapted to the hydrolysate will result in yields similar to those found for 8.4% WIS used in this study. This is also indicated by preliminary experimental results (Rudolf et al., 2005). In a near term future improvements to yeast and enzymes as well as in equipment design will make it possible to operate the SSF at a higher concentration than 12%. This will decrease the production cost further (Fig. 8). In the MVR configuration, the compressor needs electricity to reduce the energy demand in the form of steam. Thus, the outcome of this kind of evaluation will be dependent on the difference between the cost of the steam and the cost of electricity. A sensitivity analysis was conducted to estimate the effect of an increase in the cost of electricity on the ethanol production cost for the MVR cases. The ethanol production cost increases linearly as the cost of electricity increases, see Fig. 9. For 8.4% WIS in the SSF step the ethanol production cost for the AD case is the same as for the MVR case when the price of electricity is around 300 SEK/MW h. At 12% WIS breakeven occurs at a cost of around 400 SEK/MW h. Given the uncertainties in the electricity market in Sweden this result shows that no clear conclusion can be drawn as to whether the MVR configuration is more promising than the AD configuration or not. Breakeven between MVR and multiple-effect evaporation with eight effects occurs at a cost of electricity higher than 900 SEK/MW h at all levels of WIS. These results clearly show the potential for MVR as an alternative to multipleeffect evaporation. An alternative to producing the steam on-site is to locate the ethanol plant adjacent to an existing facility such as a heat and power plant or a pulp and paper mill as such plants may have a surplus of steam that can be purchased
3.4 200
400
600
800
1000
1200
Cost of electricity (SEK/MWh)
Fig. 9. Ethanol production cost as a function of the cost of electricity for the mechanical vapour recompression alternatives. Squares indicate breakeven with the corresponding (i.e. the same concentration of WIS in SSF) AD case, and triangles indicate breakeven with eight effects in the evaporation step.
at a low cost. It is therefore of interest to determine the highest acceptable cost of purchased steam, i.e. the cost that results in the same ethanol production cost as when steam is produced on-site. Thus, an evaluation was carried out in which the steam boiler, which is not needed if steam is purchased, was removed from the base case and the entire solid residue was sold at a price of 140 SEK/ MW h. Under these conditions the cost of purchased energy must be 200 SEK/MW h (corresponding to a steam cost of around 120 SEK/ton at an average energy requirement of 2.2 MJ/kg steam produced) to obtain the same overall ethanol cost as the base case. Thus, if steam can be purchased at a cost lower than 200 SEK/MW h the co-location scenario is of interest compared with a standalone facility. As stated above, the implementation of an anaerobic digester is associated with great uncertainties since experimental data are lacking. Whether it will be a reasonable alternative to evaporation will be dependent on its efficiency in terms of the removal of COD and inhibitory compounds. The energy efficiency, i.e. the amount of biogas that can be produced from the COD in the stillage is also of importance. Another sensitivity study was carried out to investigate the effect of the efficiency of the anaerobic digester in terms of methane production on the ethanol production cost. The amount of methane produced was varied while all other costs were kept constant except the income from the solid fuel co-product. According to Fig. 10 the ethanol production cost decreases with increased methane production due to a decrease in the amount of solid residue needed for steam generation. The theoretical cost reduction is 0.2 SEK/L independent of the concentration of WIS in the SSF. At zero methane production the ethanol production cost is 4.02 SEK/L at 8.4%
A. Wingren et al. / Bioresource Technology 99 (2008) 2121–2131
recycling. A method must also be found for disposing of the sludge.
Ethanol production cost (SEK/L)
4.2 8.4% WIS 10.0% WIS 12.0% WIS
4.0
References
3.8
3.6
3.4
3.2
3.0
0
200
400
600
800
1000
1200
2131
1400
Amount of methane produced (kg/h)
Fig. 10. Ethanol production cost as a function of methane production in the AD configuration for 8.4%, 10% and 12% WIS in the SSF step.
WIS in SSF. This cost is lower than that of the base case process where the residual sugar stream is concentrated and thereafter used for steam production or sold. Thus, even for a process configuration where the evaporation system is replaced with an aerobic digester alone and no methane is produced, a lower ethanol production cost than the base case would be obtained. This clearly demonstrates the high cost of evaporating such a diluted stream as the one corresponding to 8.4% WIS in the SSF step.
5. Conclusions The results of this study indicate that both anaerobic digestion and mechanical vapour recompression applied to the evaporation step are promising alternatives to traditional multiple-effect evaporation for the treatment of the stillage. These two configurations have lower energy demands and result in a lower ethanol production cost than the base case. A sensitivity study revealed that the anaerobic digester option is very promising, even if the efficiency in terms of methane production is very low. However, experimental work will be necessary to determine its performance and whether the treated water is suitable for
Aden, A., Ruth, M.F., Ibsen, K., Jechura, J., Neeves, K., Sheehan, J., Wallace, B., 2002. Lignocellulosic biomass to ethanol processing design and economics utilizing co-current dilute acid prehydrolysis and enzymatic hydrolysis for corn stover. NREL/TP-510-32438. Alkasrawi, M., Galbe, M., Zacchi, G., 2002. Recirculation of process streams in fuel ethanol production from softwood based on simultaneous saccharification and fermentation. Appl. Biochem. Biotechnol., 849–861. Alkasrawi, M., 2004. Development of Simultaneous Saccharification and Fermentation for Production of Ethanol from Softwood. Ph.D. Thesis. LUTKDH/(TKKA-1006)/1-51/(2004), Dept. of Chemical Engineering, Lund University. Bergeron, P., 1996. Environmental impacts of bioethanol. In: Wyman, C.E., (Ed.), Handbook on Bioethanol: Production and Utilization, pp. 89–103. Callander, I.J., Clark, T.A., McFarlene, P.N., 1986. Anaerobic digestion of wood ethanol stillage using upflow anaerobic sludge blanket reactor. Biotechnol. Bioeng. 30, 896–908. Gregg, D., Saddler, J.N., 1996. A techno-economic assessment of the pretreatment and fractionation steps of a biomass-to-ethanol process. Appl. Biochem. Biotechnol. (57/58), 711–727. Larsson, M., Galbe, M., Zacchi, G., 1997. Recirculation of process water in the production of ethanol from softwood. Bioresour. Technol. 60 (2), 143–151. Lynd, L.R., Elander, T.E., Wyman, C.E., 1996. Likely features and costs of mature biomass ethanol technology. Appl. Biochem. Biotechnol. (57/58), 741–761. Rudolf, A., Alkasrawi, M., Zacchi, G., Lide´n, G., 2005. A comparison between batch and fed-batch simultaneous saccharafication and fermentation of steam pretreated spruce. Enz. Microbial. Technol. 37, 195–204. Shojaosadati, S., 1996. The use of biomass and stillage recycle in conventional ethanol fermentation. J. Chem. Technol. Biotechnol. 67 (4), 326–366. St. Julian, G., Bothast, R.J., Krull, L.H., 1990. Glycerol accumulation while recycling thin stillage in corn fermentations to ethanol. J. Ind. Microbiol. 5, 391–394. von Sivers, M., Zacchi, G., 1995. A techno-economical comparison of three processes for the production of ethanol from pine. Bioresour. Technol. 51, 43–52. Wilkie, A.C., Riedesel, K.J., Owens, J.M., 2000. Stillage characterization and anaerobic treatment of ethanol stillage from conventional and cellulosic feedstocks. Biomass Bioenergy 19, 63–102. Wingren, A., So¨derstro¨m, J., Galbe, M., Zacchi, G., 2004. Process considerations and economic evaluation of two-step steam pretreatment for the production of fuel ethanol from softwood. Biotechnol. Progr. 20 (5), 1421–1429.
Bioresource Technology 99 (2008) 2121–2131
Energy considerations for a SSF-based softwood ethanol plant Anders Wingren *, Mats Galbe, Guido Zacchi Department of Chemical Engineering, Lund University, P.O. Box 124, SE-221 00 Lund, Sweden Received 8 September 2006; received in revised form 23 May 2007; accepted 24 May 2007 Available online 27 September 2007
Abstract Ethanol can be produced from softwood by steam pretreatment followed by simultaneous saccharification and fermentation (SSF). However, the final ethanol concentration in the SSF step is usually rather low (around 4 wt%) and as a result the energy demand in the downstream processing will be high. In an effort to reduce the energy consumption various alternatives for the downstream processing part of the process were evaluated from a technical-economic standpoint. With experimental data as a basis, the whole process was modelled using the commercial flowsheeting program Aspen Plus. The results were used in the subsequent economic evaluation, which was performed using Icarus process evaluator. A base case configuration, consisting of three thermally coupled distillation columns and multiple-effect evaporation was established. For a feed containing 3.5% ethanol (w/w) to the distillation step, the overall process demand for steam was estimated to be 19.0 MJ/L ethanol and the ethanol production cost 4.14 SEK/L (0.591 USD/L). Different alternatives were considered, such as integration of a stripper with the evaporation step, increasing the number of evaporation effects and the application of mechanical vapour recompression to the evaporation step. Replacement of evaporation with anaerobic digestion was also considered. Among these alternatives, evaporation using mechanical vapour recompression and the anaerobic digester alternative both resulted in significantly lower energy demand than the base case, 10.2 and 9.8 MJ/L, respectively, and productions costs of 3.82 (0.546 USD/L) and 3.84 SEK/L (0.549 USD/L). Ó 2007 Elsevier Ltd. All rights reserved. Keywords: Ethanol; Economics; Process; SSF; Energy optimization
1. Introduction We are now having to face the fact that the world’s resources of oil will eventually be depleted and that alternative fuels will have to be found. The optimal replacement fuel should be cheap and it should be possible to produce it in large quantities from a renewable source. Furthermore, the life cycle of the fuel should be environmental friendly. It would be an advantage if this fuel could also be used directly in existing motor vehicles and the same distribution system could be used as for petrol. Bioethanol exhibits almost all these desirable properties. It can be produced from lignocellulosic feedstocks, which are cheap, renewable and available in large quantities. If bioethanol *
Corresponding author. Tel.: +46 46 2228297; fax: +46 46 2224526. E-mail addresses: [email protected] (A. Wingren), Guido. [email protected] (G. Zacchi). 0960-8524/$ - see front matter Ó 2007 Elsevier Ltd. All rights reserved. doi:10.1016/j.biortech.2007.05.058
were to replace petrol the emission of greenhouse gases would be reduced by more than 85% when considering the whole fuel cycle (Bergeron, 1996). Ethanol also has the advantage of being easy to distribute. Gasoline–ethanol blends containing up to around 20% ethanol can readily be used in modern spark-ignited combustion engines, while flexible fuel vehicles can run on conventional petrol, ethanol or combinations of the two. Various studies have concluded that it is possible to produce ethanol from lignocellulosic raw materials in an economically feasible way (Aden et al., 2002; Lynd et al., 1996; von Sivers and Zacchi, 1995; Wingren et al., 2004). In Sweden softwood is considered to be the raw material with the highest potential for ethanol production. It is abundant and has a high content of carbohydrates that can be hydrolysed to fermentable sugars. In previous techno-economic evaluations of softwood as raw material, a process consisting of steam pretreatment followed by
2122
A. Wingren et al. / Bioresource Technology 99 (2008) 2121–2131
enzymatic hydrolysis and fermentation was studied (von Sivers and Zacchi, 1995; Wingren et al., 2004). In this process the lignin-rich residue is a valuable co-product that can be dried, pelletized and sold as a solid fuel. However, a major part of the solid residue is needed for the production of process steam. Compared with sugar-based or starchbased processes the feed to the distillation step in a future lignocellulosic plant is likely to be more dilute, and as a consequence, the downstream processing will be energy demanding and thus costly. To increase the income from this co-product the energy demand in the process has to be reduced. Several options are available. (i) A higher substrate load in the saccharification and fermentation steps increases the ethanol concentration in the feed to the downstream processing steps. Thus the energy demand in the distillation and evaporation steps decreases. (ii) The fresh water added in the saccharification and fermentation steps can be replaced with recirculated process streams such as the stream before distillation or the stillage stream after distillation (Alkasrawi et al., 2002). The latter is an option widely employed in sugar-based and starch-based ethanol processes (Shojaosadati, 1996; St. Julian et al., 1990). (iii) The energy demand can be reduced by energy integration where the use of secondary steam and waste heat is optimized. (iv) Another option is to optimize the energydemanding process steps directly, e.g. by employing more advanced distillation or evaporation systems or by replacing these steps with less energy-intensive unit operations. The main objectives of downstream processing are to produce fuel ethanol, to recover valuable co-products and to produce wastewater that can either be recycled in the process or easily be disposed of. Although there are several options, distillation is almost the only unit operation used for concentration of dilute ethanol streams in commercial ethanol processes. Several treatment methods are available for the stillage and an excellent review paper on this topic has been written by Wilkie et al. (2000). In sugar- and starch-based ethanol plants evaporation, anaerobic digestion, membrane filtration and direct application to land have been used as methods of stillage treatment. Combinations of these methods are also common. Regarding lignocellulosic feedstocks, only limited data regarding the treatment of the stillage have been published. Evaporation is an energy-demanding unit operation but it produces a syrup that can be combusted, and clean condensates that can be recycled or treated further before being disposed of. Alkasrawi (2004) studied the effect of recycling evaporation condensates in a simultaneous saccharification and fermentation (SSF) process, and Larsson et al. (1997) performed similar studies on a separate hydrolysis and fermentation (SHF) process. Their conclusions were that a major part of the fresh water added to the process could be replaced with condensates without affecting ethanol yield or production rate. Anaerobic digestion, followed by an aerobic step, is an interesting option since it has a low energy demand and produces methane, which can be used for steam production. Callander et al. (1986) studied
anaerobic digestion of softwood stillage with a COD and BOD content of 25 000 and 13 000 mg/L, respectively. A removal of 86% of the COD and 93% of the BOD was achieved. The effect of recycling this purified stream was not investigated. This study deals with the ‘‘ethanol-from-wood’’ process from an energy perspective. A base case is established and used for comparison with alternative process configurations in which energy-saving options, such as multiple-effect evaporation and integration of evaporation and distillation are employed. The use of mechanical vapour recompression and replacement of evaporators with an anaerobic digester have also been investigated. Since downstream processing is highly energy demanding, distillation and evaporation were studied in detail, but as the energy requirement of a plant of this kind cannot be optimized by the investigation of a single step, the fully integrated process was evaluated. 2. Process description 2.1. Pretreatment and SSF Fig. 1 shows the flowsheet for the process on which the simulations were based on. Fresh wood chips of spruce are impregnated with sulphur dioxide (2 wt% of the liquid in the raw material) and thereafter steam pretreated at 215 °C for 5 min. The slurry is flash-cooled by pressure reduction to 4 bar and then to 1 bar, after which it is fed to the SSF step. The pH-value is adjusted with lime to about 5 and the slurry diluted to a concentration of water-insoluble solids (WIS) of 8.4%. After temperature adjustment to 37 °C and addition of enzymes, corresponding to 15 filter paper units per gram WIS, and nutrients, the conversion starts and is completed in 72 h. The process is assumed to be operated batchwise with a total cycle time of 84 h. Process data for these steps are based on experimental work carried out in a process development unit (PDU) at the Department of Chemical Engineering, Lund University, Sweden. The dry raw material consists of 45.0% glucan, 12.6% mannan, 2.6% galactan, 7.1% pentosans, 28.1% lignin and the remainder is made up of acetyl groups, extractives and ash. The water content is 50%. In the pretreatment step 15% of the glucan is converted to glucose and 10% to byproducts such as 5-hydroxymethyl furfural. Of the mannan, galactan, arabinan and xylan, 70% is converted to monomeric sugars and 25% to by-products. Around 5% of the lignin is solubilized. The mass recovery, based on dry weight, in this step is 90% of the dry raw material. In the subsequent SSF step 94% of the glucan is hydrolysed. Of the glucose and mannose, 92% is converted to ethanol and the remaining fraction is converted to by-products or remains unfermented. All other sugars fed to the SSF step are unconverted. The major difference between the experimental work and the simulation is the mode of operation in the SSF step which was fed-batch in the experimental work
A. Wingren et al. / Bioresource Technology 99 (2008) 2121–2131
Wood
Feedstock handling
Pretreatment
SSF
2123
Ethanol
Distillation Stillage
Evaporation
Solid-liquid separation
Pelletizing Drying
Solid fuel
Solids
Incineration
Steam
Syrup
Other
Fig. 1. Flowsheet for the ethanol-from-wood process.
but assumed to be batch in the simulation and in the economic evaluation. Also, baker’s yeast (Saccharomyces cerevisiae) was added to a concentration of 5 g/L in the experiments. In a full-scale process it is more likely that the yeast would be produced in-house using some of the sugar available in the hydrolysate as substrate. In the simulation it was therefore assumed that the yeast is produced directly in the fermentors to a concentration of 2 g/L prior to the start of SSF. 2.2. Distillation and evaporation The flowsheet for this step is shown in Fig. 2. The base case consists of two strippers and a rectifier with top stage pressures of 3, 1.25 and 0.35 bar. The stream from the SSF step is preheated and then divided between the two strippers. Overhead vapour from the stripper at 3 bar is used as the heat source in the reboiler of the second stripper before being fed to the rectification column. Overhead
vapour from the second stripper, as well as some primary steam, is used as the heat source in the reboiler of the rectification column. The feed to the two strippers consists of the unfiltered mash containing not only water and ethanol, but also solids such as lignin and yeast. Therefore, it can be expected that the tray efficiency will be quite modest, and in this study a Murphree efficiency of 50% was used. The rectifier, operating without solids, is expected to have a higher efficiency, 75%. The recovery of ethanol was assumed to be 99.5% in all three columns and the distillate from the rectifier was assumed to consist of ethanol at 94% w/w. For the specifications given above for the rectifier the reflux ratio is dependent on the number of trays. A higher reflux ratio reduces the number of trays, makes the column diameter larger and requires a higher reboiler duty. Thus, the optimal reflux ratio, i.e. that resulting in the lowest overall cost, is a trade-off between capital cost and energy cost. Simulations were carried out to study how the number of trays
Ethanol (94% w/w)
Stripping
Evaporation
Rectification
Centrifugation
Syrup (50% DM)
From SSF
Fig. 2. Flowsheet for the distillation and the evaporation steps in the base case. Units using primary steam or secondary steam from other process steps, i.e. the dryer or the pretreatment are shaded. Dotted lines indicate vapour phase.
2124
A. Wingren et al. / Bioresource Technology 99 (2008) 2121–2131
affected the reflux ratio and the reboiler duty, and it was concluded that a reasonable number of trays was 35 in the rectifier with a reflux ratio of 2.4. The two strippers are equipped with 25 trays each. The stillage is sent to a liquid–solid separation step consisting of two decanter centrifuges connected in parallel. The concentration of WIS in the solid fraction is assumed to be 30%. In the evaporation step the liquid is concentrated to 50% dry matter (DM) and part of the condensate can be recycled in the SSF step. A multiple-effect evaporator system consisting of five effects with a forward feed arrangement was used in the base case configuration. The pressure in the first effect is 3 bar and in the last 0.2 bar. The pressures in effects 2–4 were adjusted to achieve roughly the same difference in temperature between the hot and cold side of all effects. The boiling point elevation for wood hydrolysates has been determined experimentally (Larsson et al., 1997) and was accounted for. Overall heat transfer coefficients were varied from 2000 to 600 W/m2 °C depending on the temperature and concentration of the slurry being evaporated. Outgoing condensate is used for preheating of the feed to the strippers. 2.3. Steam generation and solid fuel production The concentrated stream from the evaporator and the solid stream from the centrifuges are mixed and fed to the drying step where the material is dried to 85% DM. The equipment is assumed to be a steam dryer in which superheated steam is used as the drying medium. Primary steam is required at two pressure levels, 20 and 12 bar, and the dryer works at 4 bar, which means that secondary steam is available at a high temperature and, thus, the net energy demand per unit evaporated liquid is low. Part of the dried solid residue is sent to the boiler where fresh
steam for the process is generated. Excess solid residue is pelletized and sold as a solid fuel. 2.4. Alternative process configurations Different process configurations for the distillation and evaporation steps were investigated. The process steps on the upstream side of these steps are unaffected. The following options were considered. 2.4.1. Increasing the number of effects in the evaporator The number of effects in the evaporation step was increased from the base case value of five up to eight effects. The pressures in the first and the last effect were the same as in the base case. The highest possible number of effects is determined by the overall temperature difference over the evaporation step. 2.4.2. Integration of the stripper with evaporation A single stripper working at 1 bar is integrated in the evaporation train with five effects, see Fig. 3. The steam from one of the evaporation effects is used as the heating medium in the reboiler of the stripper. The vapour overhead from the stripper is used as the heating medium in the following effect in the evaporation step. Thus, in this case the rectifier, which is operated at atmospheric pressure, must use primary steam as its heating source. 2.4.3. Mechanical vapour recompression applied to the evaporator A large portion of the energy supplied to a traditional multiple-effect evaporator ends up as latent heat in the vapour phase leaving the last effect in the evaporator. This vapour is normally condensed using cooling water. Another option is to compress the vapour, thereby raising the
Ethanol (94% w/w)
Stripping Rectification To drying
Evaporation
From SSF
Fig. 3. Flowsheet for the stripper integrated with the evaporation unit.
A. Wingren et al. / Bioresource Technology 99 (2008) 2121–2131
temperature to a level at which the latent heat will be available. The vapour can then be used as a heating medium to replace most of the primary steam. When compression is carried out by a mechanical compressor the process is referred to as mechanical vapour recompression (MVR). An electrical motor or a steam turbine provides the thrust to the compressor. Another option is to thermally recompress the vapour by use of a steam ejector. This process is called thermal vapour recompression. This study was limited to the use of MVR with an electrical motor. The economy is dependent on the power demand of the compressor and this is governed by the vol-
2125
umetric flow rate of the vapour and the pressure increase over the compressor. The pressure increase governs the driving force in the evaporator (i.e. the temperature difference), which affects the area needed for heat transfer. Thus, as the electricity consumption increases, the evaporator area decreases. The purpose of this study was not to fully optimize this system, but to investigate the potential of MVR in an ethanol plant. Compression from 3 to 4 bar in a single step (two effects in parallel) was assumed providing a temperature difference of 10 °C, see Fig. 4. In the MVR unit the slurry is concentrated to 20% and thereafter to 50% in a small 4-effect evaporator. The reason for not
Ethanol (94% w/w)
Stripping Compresser
Rectification
To drying
Evaporation
From SSF
Fig. 4. Mechanical vapour recompression applied to the evaporation step.
Ethanol (94% w/w)
Rectification
Stripping
To drying
From SSF
Biogas to boiler
Anaerobic digestion
Aerobic step
Purified water
Fig. 5. Evaporation replaced by anaerobic digestion followed by aerobic digestion.
2126
A. Wingren et al. / Bioresource Technology 99 (2008) 2121–2131
concentrating the slurry to 50% in the MVR unit directly is the boiling point elevation, which is 7 °C at 50% but only 1 °C at 20% for this kind of liquid (Larsson et al., 1997). 2.4.4. Anaerobic digestion The organic material in the stillage is converted through anaerobic digestion (AD) to biogas consisting of methane and carbon dioxide, see Fig. 5. The performance of such a system is dependent on a number of parameters such as the composition of the feed, residence time, temperature, etc. Since very limited data regarding the performance of this kind of system have been published only a rough estimate could be made based on data from PURAC AB, Lund, Sweden. The liquid part of the stillage is cooled to the operating temperature of the anaerobic digester (55 °C) by heat exchange with the feed to the distillation step. The production of methane is assumed to be 0.35 m3/kg COD consumed and in the anaerobic step around 50% of the COD is removed. In the subsequent aerobic step most of the remaining COD is removed and this water can then be either recycled to the SSF step or disposed of. The residence time in the anaerobic step is 20 days and in the aerobic step 4 days. Nitrogen and phosphor are added at a rate corresponding to 26 kg ammonia/h and 23 kg phosphoric acid/h. The biogas produced is sent to the steam boiler. Sludge will also be produced, especially in the aerobic step. However, the amount and composition are associated with great uncertainties. Sludge can be dewatered and thereafter landfilled or incinerated. These aspects were not studied in detail in this study, and it was assumed that the contribution of sludge to the overall production cost is minor and was thus neglected. The process configurations studied were all based on the same fundamental process design. Primary steam was assumed to be available at 25, 12 and 6.9 bar and secondary steam was used to replace fresh steam when possible. The plant considered is a grass-root plant, designed to process 200 000 tonnes of raw material annually on a DM basis. Annual on-stream time was assumed to be 8000 h. 3. Methods The process simulations were carried out using Aspen Plus version 11.1 from Aspen Technology. All process steps described above were simulated using built-in modules for reactors, heat exchangers, pumps, etc. Distillation columns were simulated rigorously, i.e. using the RADFRAC module. Results from the process simulations were used in the sizing of the process equipment. Capital costs for standard equipment such as heat exchangers, columns, evaporators, vessels and pumps were estimated using Icarus process evaluator (IPE), version 12.0 from Aspen Technology. Auxiliary equipment such as piping, electrical equipment and instrumentation, as well as indirect costs arising from engineering, freight, construction indirects and contingency, were also estimated using IPE. The equipment cost estimated by IPE was compared with quo-
tation costs obtained from Swedish suppliers of the same kind of equipment and showed good agreement. The pretreatment reactor was assumed to work like a StakeTech reactor and a quotation provided from Stake Technology was used in the cost estimation. Cost data for the pellet machine, anaerobic digester, compressor for the mechanical vapour recompression and the dryer were also obtained from vendors, and data for the steam boiler were obtained from the literature. Working capital was accounted for as the cost of raw material and chemicals for 10 days production, the value of final products for 30 days production, 30 days of operating expenses and accounts receivable to allow a 30-day payment credit for customers. The annual capital cost was determined by multiplying the fixed capital investment by an annuity factor of 0.103 corresponding to a 15-year pay-off time and an interest rate of 6%. This should be regarded as a favourable bank loan. A higher interest rate and/or a shorter pay-off time would result in a higher annuity factor and thus a higher annual capital cost. The cost of the raw material was assumed to be 90 SEK/MW h, while the income from the solid fuel coproduct was 140 SEK/MW h. Costs used in this evaluation are summarized in Table 1. The purpose of the process simulations and the economic evaluation was to compare different process configurations on an equal basis and not to obtain an absolute ethanol production cost. Table 1 Costs used in the economic evaluation Type
Cost
Unit
Raw material Wooda
90
SEK/MW h
Chemicals SO2 CaO Defoamer NaOH (50%) (NH4)2HPO4 MgSO4 Æ 7H2O NH3 (29%) H3PO4 (50%) Enzymes
1.5 0.7 20 2.0 1.5 4.4 2.6 5.0 19
SEK/kg SEK/kg SEK/kg SEK/kg SEK/kg SEK/kg SEK/kg SEK/kg SEK/106 FPU
By-product income Solid fuelb Carbon dioxide
140 0.03
SEK/MW h SEK/kg
Utilities Electricity Cooling water Process water
250 0.14 1.40
SEK/MW h SEK/m3 SEK/m3
Other costs Labour
500 000
Insurance Maintenance
1 2
SEK/employee year (25 employees assumed) % of fixed capital per year % of fixed capital per year
Capital costs Annuity factor
0.103
a b
0.42 SEK/kg DM. 0.79 SEK/kg DM.
A. Wingren et al. / Bioresource Technology 99 (2008) 2121–2131
2127
steam pretreatment of hardwood using roughly the same process conditions as in this study. They concluded that approximately 50% of the sulphur added in the impregnation step remained in the slurry following steam pretreament. Only a small fraction (7% of the added sulphur) was bound to the water-insoluble fraction, and the sulphur in the liquid phase was mainly in a non-volatile form such as lignosulphonates. This result clearly indicates that most of the sulphur added in the pretreatment step in the evaluated plant will be found in the syrup stream leaving the evaporation unit and not in the fibrous stream. Since the excess solid residue is intended to be sold as a solid fuel, a low sulphur content, a high heating value and a low ash content are desirable. The syrup can be combusted directly instead of mixing it with the fibrous stream prior to drying. Part of the dried fibrous stream can also be incinerated as well if necessary. Further research is needed in this area. Table 2 lists the energy demand in the different process steps for the various alternatives. In the base case configuration the process steps with the highest net energy demand are distillation and evaporation. These two steps require 29.6 MW steam. Pretreatment and drying require high pressure steam but also produce significant amounts of secondary steam (16.3 MW), which can replace some of the primary steam needed in the distillation and evaporation steps. The total energy demand in the form of primary (boiler-generated) steam is thus 38.8 MW or 19.0 MJ/L ethanol. This steam is generated by incinerating 61% of the 12 000 kg/DM h of solids from the drying step thus leaving 4650 kg DM/h for sale. Fig. 6 shows an energy flow diagram of the base case configuration. The energy bound in the raw material is divided mainly between ethanol and solid fuel. Of the 143 MW originating from the raw material, 47.9 MW or 34% ends up as ethanol. The excess solid fuel represents 28.9 MW or 20%. Thus,
4. Results and discussion With experimental data as the basis the process was simulated in Aspen Plus to determine unknown process parameters such as temperatures, size and composition of process streams, as well as utility requirements in the different process steps. For a process with 8.4% WIS in the SSF step the concentration of ethanol in the feed to the distillation step is 3.5% (w/w) and the concentration of DM in the stream being fed to the evaporation or anaerobic digestion step is 3.3%. Ethanol is produced at a rate of 5813 kg/h, which corresponds to an overall ethanol yield of 71% of the theoretical based on the hexoses fermentable by baker’s yeast (i.e. glucan and mannan) in the raw material. The total amount of solid residue before the steam generation step, 12 000 kg DM/h, is the same for all process configurations except the AD case where the stillage stream, having a flow of dry matter of around 4200 kg/h, is fed to the anaerobic digester. The flow of solid residue before steam generation is thus lower, 7800 kg DM/h. The methane production in the AD case was estimated to be 667 kg/h. There is a significant difference between the solid residue in the AD case and in the other process configurations. In the AD configuration, the residue consists of around 80% lignin and the remaining fraction is made up of nondegraded polysaccharides, yeast, etc. In the other cases the syrup from the evaporator has been mixed with the fibrous stream prior to drying. Since the syrup consists of non-volatile solubles such as glycerol, dissolved lignin, sugars, ash, inorganic compounds, etc., the lignin content of the solid residues decreases and is estimated to be only 55%. This solid stream will have a lower heating value than the AD residue and possibly also have a lower quality as a fuel due to its higher ash content. The syrup may also contain more sulphur than the fibrous stream. (Gregg and Saddler (1996) calculated a sulphur balance for SO2-catalysed
Table 2 Overall heat duties (MW) for various process configurations, as well as the contribution from the individual process steps to the overall heat duty BC
Steam required (MW) Pretreatment Distillation Evaporation Drying
MEE8
Intstripp
MVR
HP
LP
HP
LP
HP
LP
HP
10.0
4.6 15.8b 13.8
10.0
4.6 15.9 7.7
10.0
4.6 9.7 18.4
10.0
10.9
Steam generated (MW) Pretreatment Drying
10.8 6.2 10.1
Net energy demanda
20.9
Thermal energy boiler (MW)
38.8
17.9
11.3 6.2 10.1
20.9 32.9
12.0
37.3
LP 4.6 8.4 2.5
10.9 6.2 10.5
21.3
AD
16.0
HP
LP
10.0
4.6 11.3
8.6 6.2 10.1
20.9 20.9
0.9
6.2 8.4 18.6
1.4
20.0
Data are presented for 8.4% WIS in the SSF step. BC = base case, MEE8 = multiple-effect evaporation with eight effects, Intstripp = stripper integrated with evaporator, MVR = mechanical vapour recompression, AD = anaerobic digestion. HP: High pressure steam 25 and 12 bar, LP: low pressure steam 4 and 1 bar. a Difference between steam required and steam generated. b Of which 7.4 MW is needed for preheating of the feed.
2128
A. Wingren et al. / Bioresource Technology 99 (2008) 2121–2131 Solids 55.9 (54.3)
Losses and cooling 3.6
Pretreatment
Hydrolysate 129 (126)
SSF
Broth 129 (126)
Flash condensate18.3 (17.3)
Wood 143 (143)
Excess solid fuel 29.3 (28.9) Drying
37.3 (22.1)
Fuel 46.4 (45.8)
Stillage 93.9 (76.5)
Ethanol 47.9 (47.9)
Solid sep.
Distillation
Condensates 5.2 (1.2)
Preheating 6.3
Condenser 18.4
Evaporation
Solid fuel 20.9 (20.7)
Condenser 5.7
Incineration
Losses 7.7 (1.8)
LP steam 15.8
HP steam 10.0 LP steam 4.6 LP steam 6.2
Steam distribution
LP steam 13.8 HP steam 10.9 LP steam 10.1 Steam 38.8
Fig. 6. Energy flows (in MW) in the base case configuration at 8.4% WIS in the SSF step. Raw material intake is 25 000 kg DM/h. The values given are energy flows (based on heat of combustion) calculated relative to a reference state of 25 °C and water in liquid phase. Values in brackets are without water. Heat flows, representing the net flow of energy, are shown as dotted lines. Balances may be inconsistent due to rounding-off errors and due to the exclusion of minor streams.
the overall energy efficiency is 54%. However, the process also requires 5.5 MW electricity. If this figure is included, assuming an electricity-to-fuel ratio of 0.3, the energy efficiency becomes 48%. Thirty-two percent of the energy in the raw material is used for steam production, while 12% ends up in the flash steam from the pretreatment step. The latter figure is associated with large uncertainties due to the difficulty in determining the composition of the flash steam. In the PDU, around 90% of the dry material is recovered as dry matter in the slurry following steam pretreatment. The fraction of the material not recovered as dry matter is likely to be volatile substances formed during pretreatment as with extractives from the wood. The exact composition is, however, unknown. All the alternative process configurations result in a lower energy demand than the base case. The energy demand is reduced as the number of effects in the evaporation step is increased. When eight effects are used the energy demand is reduced to 32.9 MW, while the case with an integrated stripper results in a minor reduction in energy demand. The two remaining alternatives, the MVR case and the AD case, both reduce the overall energy demand significantly. The use of MVR results in an overall energy demand of 20.9 MW and in the anaerobic digester case 20.0 MW. The electricity needed for the compressor in the former alternative is 2.3 MW. The energy demand in the distillation step is lower for the MVR and AD cases than for the base case despite the fact that the same process configuration is used in the distillation step. This is explained by the fact that these two steps are integrated, and if the process conditions are changed in one of the steps this will affect the other. In the case employing MVR the condensate stream, having a temperature of
144 °C, can provide all the energy needed for preheating of the feed to the strippers. When an anaerobic digester replaces the evaporation step, the feed to the strippers is preheated by the stillage, which can also supply essentially all the energy needed for preheating. In fact, the MVR and anaerobic digestion cases are so efficient that all the secondary low-pressure steam generated in the dryer and the pretreatment step cannot be utilized in the distillation and the evaporation steps. However, 4.6 MW can be used to preheat the wood in the pretreatment step to 120 °C. Despite this, the mechanical vapour recompression case has a small surplus of secondary steam corresponding to 0.9 MW, which is condensed by cooling water. The evaluated process configurations are summarized in Table 3. An overall capital investment of 889 million SEK was estimated for the base case process and the ethanol production cost was estimated to be 4.14 SEK/L (100% ethanol). The cost of the raw material corresponds to 1.41 SEK/L and the income from the solid fuel co-product is 0.50 SEK/L. The ethanol production cost decreases as the number of effects in the evaporation step is increased. However, the difference in cost between seven effects and eight effects is minor (the latter being 4.06 SEK/L), see Table 3. Increasing the number of effects further would thus not result in a lower ethanol cost but would result in a very low driving force (temperature difference) in the individual effects. Integration of the stripper in the evaporation step does not result in any major cost reductions. The drawback with this configuration is the temperature drop over the stripper, which reduces the available temperature difference over the evaporation step. The result is increased heat transfer area in the evaporators compared with the base case.
A. Wingren et al. / Bioresource Technology 99 (2008) 2121–2131
2129
Table 3 Summary of energy demand and costs for the process configurations evaluated with 8.4% WIS in the SSF step
Heat requirement (MJ/L ethanol) Total capital investment (MSEK)
BC
MEE6
MEE7
MEE8
Intstripp
MVR
AD
19.0 889
17.7 896
16.8 903
16.1 912
18.3 895
10.2 890
9.8 853
Costs (SEK/L ethanol) Raw material Solid fuel Capital Electricity Chemicals Enzymes Other
1.41 0.50 1.51 0.19 0.26 0.60 0.67
1.41 0.55 1.52 0.18 0.26 0.60 0.67
1.41 0.59 1.53 0.18 0.26 0.60 0.68
1.41 0.62 1.55 0.18 0.26 0.60 0.68
1.41 0.53 1.52 0.19 0.26 0.60 0.68
1.41 0.86 1.51 0.26 0.26 0.60 0.65
1.41 0.70 1.45 0.18 0.28 0.60 0.63
Total production cost (SEK/L)
4.14
4.10
4.07
4.06
4.12
3.82
3.84
Total production cost (USD/L)
0.591
0.586
0.581
0.580
0.589
0.546
0.549
BC = base case, MEE8 = multiple-effect evaporation with eight effects, Intstripp = stripper integrated with evaporator, MVR = mechanical vapour recompression, AD = anaerobic digestion.
4.5
Ethanol production cost (SEK/L)
The MVR case results in an overall ethanol production cost of 3.82 SEK/L. The income from the solid fuel coproduct is 0.86 SEK/L. Although the cost of the evaporation system increases, the total capital investment increases only slightly due to the cost reductions in the steam generation step. The cost of the electricity is higher in this case due to the power requirement of the compressor. Almost as low production cost, 3.84 SEK/L, is achieved with the process alternative using AD. This results in a lower capital cost as well as a higher income from the solid fuel co-product than the base case. An analysis was carried out to investigate how sensitive the ethanol production cost is to variations in the selling price of the solid fuel co-product. The price was varied from 50 SEK/MW h to 250 SEK/MW h and the results are shown in Fig. 7. The largest cost reduction is seen for the MVR case since this configuration results in the highest production of solid fuel, and is thus more sensitive to price changes. For the same reason, the base case configuration has the lowest sensitivity. At 200 SEK/MW h the cost difference between the case with the highest production cost (base case) and the case with the lowest production cost (MVR), is 0.74 SEK/L ethanol. However, if the selling price is 50 SEK/MW h this difference is only 0.17 SEK/L. When the process configurations are evaluated assuming a higher concentration of WIS in the SSF step, the production cost is reduced for all alternatives, see Fig. 8. An increased concentration of WIS results in increased ethanol concentration in the feed to the distillation step. At 10% and 12% WIS the ethanol concentration is 4.2% and 5.1%, respectively. The concentration of DM in the feed to the evaporator and the anaerobic digester also increases, being 4.0% and 4.9% for 10% and 12% WIS in SSF, respectively. The base case and the case with eight effects show a larger decrease in ethanol production cost than the AD and MVR configurations. This can be explained by the higher energy demand in the multiple-effect evaporation cases. The energy demand in downstream processing decreases when the concentration of WIS is increased in the SSF step,
BC MEE8 MVR AD
4.0
3.5
3.0 0
50
100
150
200
250
300
Price of solid fuel (SEK/MWh)
Fig. 7. Ethanol cost as a function of the price of the solid fuel co-product. Data for 8.4% WIS in the SSF step. BC = base case, MEE8 = multipleeffect evaporation with eight effects, MVR = mechanical vapour recompression, AD = anaerobic digestion.
and this effect is more noticeable for process configurations having a high energy demand than for energy-optimized systems. The ethanol production cost at 12% WIS is 3.73 SEK/L for the base case configuration and the lowest cost at this WIS concentration was obtained for the MVR configuration, being 3.57 SEK/L. As stated above, the evaluation is based on data obtained from experiments carried out at 8.4% WIS (fedbatch) in SSF, and the evaluation of the 10% and 12% WIS cases is based on the assumption that the yield, as well as the productivity, will be the same as in the 8.4% case. S. cerevisiae is a robust yeast that can tolerate ethanol concentrations well above 5% (corresponding to 12% WIS in SSF). However, wood hydrolysates contain other compounds which may be toxic to the yeast, as well as to the enzymes, and thus the ethanol yield as well as the
2130
A. Wingren et al. / Bioresource Technology 99 (2008) 2121–2131 4.8
4.2 BC MEE8 AD MVR
4.6
Ethanol production cost (SEK/L)
Ethanol production cost (SEK/L)
4.1
4.0
3.9
3.8
3.7
3.6
3.5 8
8.4% WIS 10% WIS 12% WIS
4.4
4.2
4.0
3.8
3.6
9
10
11
12
13
% WIS in SSF
Fig. 8. Ethanol production cost as a function of WIS for the process configurations evaluated. BC = base case, MEE8 = multiple-effect evaporation with eight effects, MVR = mechanical vapour recompression, AD = anaerobic digestion.
productivity may be reduced as the concentration of WIS increases. It is the authors’ belief, however, that an optimized fed-batch process at 12% WIS with a yeast adapted to the hydrolysate will result in yields similar to those found for 8.4% WIS used in this study. This is also indicated by preliminary experimental results (Rudolf et al., 2005). In a near term future improvements to yeast and enzymes as well as in equipment design will make it possible to operate the SSF at a higher concentration than 12%. This will decrease the production cost further (Fig. 8). In the MVR configuration, the compressor needs electricity to reduce the energy demand in the form of steam. Thus, the outcome of this kind of evaluation will be dependent on the difference between the cost of the steam and the cost of electricity. A sensitivity analysis was conducted to estimate the effect of an increase in the cost of electricity on the ethanol production cost for the MVR cases. The ethanol production cost increases linearly as the cost of electricity increases, see Fig. 9. For 8.4% WIS in the SSF step the ethanol production cost for the AD case is the same as for the MVR case when the price of electricity is around 300 SEK/MW h. At 12% WIS breakeven occurs at a cost of around 400 SEK/MW h. Given the uncertainties in the electricity market in Sweden this result shows that no clear conclusion can be drawn as to whether the MVR configuration is more promising than the AD configuration or not. Breakeven between MVR and multiple-effect evaporation with eight effects occurs at a cost of electricity higher than 900 SEK/MW h at all levels of WIS. These results clearly show the potential for MVR as an alternative to multipleeffect evaporation. An alternative to producing the steam on-site is to locate the ethanol plant adjacent to an existing facility such as a heat and power plant or a pulp and paper mill as such plants may have a surplus of steam that can be purchased
3.4 200
400
600
800
1000
1200
Cost of electricity (SEK/MWh)
Fig. 9. Ethanol production cost as a function of the cost of electricity for the mechanical vapour recompression alternatives. Squares indicate breakeven with the corresponding (i.e. the same concentration of WIS in SSF) AD case, and triangles indicate breakeven with eight effects in the evaporation step.
at a low cost. It is therefore of interest to determine the highest acceptable cost of purchased steam, i.e. the cost that results in the same ethanol production cost as when steam is produced on-site. Thus, an evaluation was carried out in which the steam boiler, which is not needed if steam is purchased, was removed from the base case and the entire solid residue was sold at a price of 140 SEK/ MW h. Under these conditions the cost of purchased energy must be 200 SEK/MW h (corresponding to a steam cost of around 120 SEK/ton at an average energy requirement of 2.2 MJ/kg steam produced) to obtain the same overall ethanol cost as the base case. Thus, if steam can be purchased at a cost lower than 200 SEK/MW h the co-location scenario is of interest compared with a standalone facility. As stated above, the implementation of an anaerobic digester is associated with great uncertainties since experimental data are lacking. Whether it will be a reasonable alternative to evaporation will be dependent on its efficiency in terms of the removal of COD and inhibitory compounds. The energy efficiency, i.e. the amount of biogas that can be produced from the COD in the stillage is also of importance. Another sensitivity study was carried out to investigate the effect of the efficiency of the anaerobic digester in terms of methane production on the ethanol production cost. The amount of methane produced was varied while all other costs were kept constant except the income from the solid fuel co-product. According to Fig. 10 the ethanol production cost decreases with increased methane production due to a decrease in the amount of solid residue needed for steam generation. The theoretical cost reduction is 0.2 SEK/L independent of the concentration of WIS in the SSF. At zero methane production the ethanol production cost is 4.02 SEK/L at 8.4%
A. Wingren et al. / Bioresource Technology 99 (2008) 2121–2131
recycling. A method must also be found for disposing of the sludge.
Ethanol production cost (SEK/L)
4.2 8.4% WIS 10.0% WIS 12.0% WIS
4.0
References
3.8
3.6
3.4
3.2
3.0
0
200
400
600
800
1000
1200
2131
1400
Amount of methane produced (kg/h)
Fig. 10. Ethanol production cost as a function of methane production in the AD configuration for 8.4%, 10% and 12% WIS in the SSF step.
WIS in SSF. This cost is lower than that of the base case process where the residual sugar stream is concentrated and thereafter used for steam production or sold. Thus, even for a process configuration where the evaporation system is replaced with an aerobic digester alone and no methane is produced, a lower ethanol production cost than the base case would be obtained. This clearly demonstrates the high cost of evaporating such a diluted stream as the one corresponding to 8.4% WIS in the SSF step.
5. Conclusions The results of this study indicate that both anaerobic digestion and mechanical vapour recompression applied to the evaporation step are promising alternatives to traditional multiple-effect evaporation for the treatment of the stillage. These two configurations have lower energy demands and result in a lower ethanol production cost than the base case. A sensitivity study revealed that the anaerobic digester option is very promising, even if the efficiency in terms of methane production is very low. However, experimental work will be necessary to determine its performance and whether the treated water is suitable for
Aden, A., Ruth, M.F., Ibsen, K., Jechura, J., Neeves, K., Sheehan, J., Wallace, B., 2002. Lignocellulosic biomass to ethanol processing design and economics utilizing co-current dilute acid prehydrolysis and enzymatic hydrolysis for corn stover. NREL/TP-510-32438. Alkasrawi, M., Galbe, M., Zacchi, G., 2002. Recirculation of process streams in fuel ethanol production from softwood based on simultaneous saccharification and fermentation. Appl. Biochem. Biotechnol., 849–861. Alkasrawi, M., 2004. Development of Simultaneous Saccharification and Fermentation for Production of Ethanol from Softwood. Ph.D. Thesis. LUTKDH/(TKKA-1006)/1-51/(2004), Dept. of Chemical Engineering, Lund University. Bergeron, P., 1996. Environmental impacts of bioethanol. In: Wyman, C.E., (Ed.), Handbook on Bioethanol: Production and Utilization, pp. 89–103. Callander, I.J., Clark, T.A., McFarlene, P.N., 1986. Anaerobic digestion of wood ethanol stillage using upflow anaerobic sludge blanket reactor. Biotechnol. Bioeng. 30, 896–908. Gregg, D., Saddler, J.N., 1996. A techno-economic assessment of the pretreatment and fractionation steps of a biomass-to-ethanol process. Appl. Biochem. Biotechnol. (57/58), 711–727. Larsson, M., Galbe, M., Zacchi, G., 1997. Recirculation of process water in the production of ethanol from softwood. Bioresour. Technol. 60 (2), 143–151. Lynd, L.R., Elander, T.E., Wyman, C.E., 1996. Likely features and costs of mature biomass ethanol technology. Appl. Biochem. Biotechnol. (57/58), 741–761. Rudolf, A., Alkasrawi, M., Zacchi, G., Lide´n, G., 2005. A comparison between batch and fed-batch simultaneous saccharafication and fermentation of steam pretreated spruce. Enz. Microbial. Technol. 37, 195–204. Shojaosadati, S., 1996. The use of biomass and stillage recycle in conventional ethanol fermentation. J. Chem. Technol. Biotechnol. 67 (4), 326–366. St. Julian, G., Bothast, R.J., Krull, L.H., 1990. Glycerol accumulation while recycling thin stillage in corn fermentations to ethanol. J. Ind. Microbiol. 5, 391–394. von Sivers, M., Zacchi, G., 1995. A techno-economical comparison of three processes for the production of ethanol from pine. Bioresour. Technol. 51, 43–52. Wilkie, A.C., Riedesel, K.J., Owens, J.M., 2000. Stillage characterization and anaerobic treatment of ethanol stillage from conventional and cellulosic feedstocks. Biomass Bioenergy 19, 63–102. Wingren, A., So¨derstro¨m, J., Galbe, M., Zacchi, G., 2004. Process considerations and economic evaluation of two-step steam pretreatment for the production of fuel ethanol from softwood. Biotechnol. Progr. 20 (5), 1421–1429.