A co-solvent hydrolysis strategy for the production of biofuels: process synthesis and technoeconomic analysis

Abstract

We develop an integrated strategy for the production of ethanol from lignocellulosic biomass. Cellulose and hemicellulose fractions are first hydrolyzed into sugars using a mixture of γ-valerolactone (GVL), water, and toluene as a solvent containing dilute sulfuric acid as a catalyst, and the sugars are then co-fermented into ethanol over engineered yeast strains. Separation subsystems are designed to effectively recover GVL and toluene for reuse in biomass hydrolysis and to recover lignin and humins for heat and power generation. We also develop an alternative process, in which we recover sugars and GVL from the residual biomass. To minimize utility requirements, we conduct heat integration, which allows us to meet all heating requirements using biomass residues. Finally, we perform a range of system-level analyses to identify the major cost and technological drivers. The proposed strategy is shown to be cost-competitive with other strategies.

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1. Introduction

Lignocellulosic biomass is a promising source for sustainable biofuel production to satisfy future energy demands while mitigating greenhouse gas emissions.^1,2 Depending on the initial treatment of biomass, the processes for upgrading biomass to fuels and chemicals can be classified into four main groups:³ (1) gasification to produce a syngas intermediate,^4–9 (2) thermochemical liquefaction and/or pyrolysis for bio-oil production,^10–17 (3) hydrolysis for sugar intermediate production,^18–26 and (4) catalytic conversion to produce levulinic acid (LA) and furfural intermediates.^27–33 Among these alternatives, biochemical hydrolysis followed by fermentation or dehydration has been investigated extensively.^34–39

Han et al. (2015)⁴⁰ proposed a nonenzymatic ethanol production strategy that employs chemical hydrolysis using γ-valerolactone (GVL) and water as a solvent containing a dilute acid catalyst,²⁰ which has a number of advantages over conventional catalytic or enzymatic ethanol production: (1) it uses GVL, which is a biomass-derived renewable solvent,^41,42 as a solvent for hydrolysis instead of expensive enzymes; (2) it produces sugars in high yields, leading to high production of ethanol; and (3) it offers flexibility because it leads to sugar intermediates which are compatible with biological or chemical upgrading strategies for a range of high value materials. They reported that their process can produce ethanol more economically than other processes (about $0.14–$0.44 per GGE of difference), such as catalytic alkenes²⁹ and enzymatic ethanol production.³⁶

Despite its advantages, the process reported by Han et al. (2015)⁴⁰ has shortcomings mostly related to separations. First, a high-pressure CO₂-based extraction system, composed of multiple compressors, is employed for the recovery of sugars from the GVL/water solution; it requires about 20% of the total installed equipment cost and 10.3 MW of electricity (for 2000 tons per day of biomass treatment). Second, an energy-intensive GVL evaporator (requires 77.8 MW heating for 2000 tons per day of biomass treatment) is used for the separation of lignin and humins from GVL.

Recently, Motagamwala et al. (2016)⁴³ have developed a new chemical hydrolysis technology that utilizes a mixture of GVL, water, and an organic co-solvent as a solvent containing dilute sulfuric acid as a catalyst. They reported three important features, which can potentially reduce separation costs while retaining the same advantages as the hydrolysis reported by Luterbacher et al. (2014):²⁰ (1) sugar could be produced in high yields, comparable to the case where we only use a mixture of GVL and water as a solvent containing a dilute sulfuric acid catalyst;²⁰ (2) an aqueous phase containing sugars was rendered immiscible with an organic phase by decreasing the temperature, while most of the GVL partitions into the organic phase; and (3) lignin and humins precipitate in the presence of the co-solvent. Among the co-solvents reported by Motagamwala et al. (2016),⁴³ toluene was of particular interest because it has the highest extraction efficiency and a moderate environmental impact.

A potential bottleneck for the practical implementation of the technology described above is the downstream separation sections necessary to recover (1) GVL and toluene for their reuse in biomass deconstruction and (2) sugars to be fermented to ethanol. Accordingly, the purpose of this study is to synthesize a new nonenzymatic ethanol production strategy that addresses these shortcomings.

In the next section, we briefly review the technologies used in our new strategy. Then, in section 3, we propose an integrated process as well as an alternative configuration. In section 4, we present the results of the technoeconomic analysis and compare the proposed strategy with the GVL/water-based nonenzymatic ethanol production⁴⁰ and enzymatic/biochemical ethanol production.³⁵ Finally, we discuss different excess biomass residue handling strategies and perform sensitivity analyses.

2. Technology review

The strategy proposed in this study integrates a catalytic conversion section for the simultaneous conversion of hemicellulose and cellulose to soluble sugars with sections for the separation of (1) toluene (to be recycled) from an aqueous phase containing lignin, humins, sugars, and GVL, (2) lignin and humins (to be used for heat and electricity generation) from sugars and GVL, and (3) sugars (to be fermented to ethanol) from GVL (to be recycled).

2.1. Nonenzymatic sugar production using toluene as a co-solvent

Motagamwala et al. (2016)⁴³ proposed a two-stage reaction system for the chemical hydrolysis of hemicellulose and cellulose to C₅ and C₆ sugars, using dilute sulfuric acid (SA) as a catalyst and biomass-derived GVL and toluene as solvents, as shown in Fig. 1. Biomass is first treated with a GVL–water (4 : 1 wt%) solvent containing 0.15 M SA for 1 h under mild conditions (390 K, 20.4 atm) with a solvent : biomass mass ratio of 4 (R-1 in Fig. 1). In the first reactor, hemicellulose is converted to C₅ sugars and furfural in 60.2% and 24.0% molar yields, respectively (reactions 1 and 2 in Table 1), while cellulose is deconstructed to C₆ sugars, LA, and 5-hydroxymethylfurfural (HMF) in 10.9%, 3.2%, and 1.4% molar yields, respectively (reactions 4, 5, and 6 in Table 1). Then, the effluent from the first reactor (stream 3) is filtered (S-1) to yield solid and liquid streams. The solid stream (stream 5), containing hemicellulose and cellulose, which are unreacted in the first reaction stage, is transferred to the second reactor for further treatment. The liquid mixture (stream 4) containing xylose, lignin, and furfural, which are soluble in GVL and water, is sent to the holding tank (V-1) where it is mixed with the hydrolyzed slurry (stream 7) discharged from the second reaction stage prior to being routed to the subsequent separation section for the recovery of sugars and biomass residues.

Fig. 1 Process flow diagram for the two-stage reaction system.

Table 1 Summary of reaction conditions and yields for the conversion of cellulose and hemicellulose

	After 1st stage	After 2nd stage
Reaction conditions
GVL : water : toluene mass ratio	4 : 1 : 0	3 : 1 : 1
Solvent to biomass mass ratio	4	16
Sulfuric acid concentration (mM)	150	5
Temperature (K)	390	430–490
Pressure (atm)	20.4	20.4
Yields (%)
(1) Hemicellulose + H2O → xylose	60.2	64.5
(2) Hemicellulose → furfural + 2H2O	24.0	30.1
(3) Hemicellulose → humins	4.8	5.4
(4) Cellulose + H2O → glucose	10.9	69.6
(5) Cellulose → LA + FA	3.2	10.4
(6) Cellulose → HMF + 2H2O	1.4	4.6
(7) Cellulose → humins	2.8	15.4

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The remaining biomass, primarily composed of cellulose and some hemicellulose, is then treated with a GVL–water–toluene (3 : 1 : 1 mass ratio) solution containing 5 mM SA for 0.67 h at 430–490 K and 20.4 atm, with a solvent : biomass mass ratio of 16 (R-2 in Fig. 1). The second stage is designed as a vertical counter-current reactor, in which biomass conversion takes place as the solids and the liquid flow in opposite directions with a progressive temperature increase for the solids. The hot liquid (490 K) is driven down the reactor and it cools to 430 K, while the cold solids are progressively heated from 430 to 490 K with the solids moving upward. For more details on a counter-current reactor, readers are referred to Song and Lee (1982)⁴⁴ and Lee et al. (1999).⁴⁵ Hemicellulose is converted to C₅ sugars and furfural in 38.8% and 55.9% molar yields, respectively, whereas cellulose is converted to C₆ sugars, LA, and HMF in 71.9%, 8.8%, and 3.9% molar yields, respectively. The remaining hemicellulose and cellulose are degraded to form humins in 5.4% and 15.4% molar yields, respectively. Based on the liquid collected from both the first and second reactors, soluble C₅ and C₆ sugars are produced in 64.5% and 69.6% molar yields, respectively (reactions 1 and 4 in Table 1). Soluble degradation products, such as furfural, LA, and HMF, are also formed in 30.1% 10.4%, and 4.6% molar yields, respectively (reactions 2, 5, and 6 in Table 1), while an insoluble degradation product (humins) is produced in 20.7% molar yield (reactions 3 and 7 in Table 1).

2.2. Separations

Motagamwala et al. (2016)⁴³ showed that the aqueous phase containing sugars was rendered immiscible with the organic phase (toluene) by decreasing the temperature below 425 K. This phase separation leads to an increase in the total soluble sugar concentration in the aqueous phase (about 94.5% of the sugars remain in the aqueous phase), while approximately two thirds of the GVL partitions into the organic phase. This separation leads to a reduction in the size of the downstream processing units and allows the recycling of the organic phase containing GVL back to the biomass deconstruction reactors using a simple liquid–liquid separator. Toluene also renders lignin and humins insoluble in the two liquid phases. Thus, biomass residues can be easily separated from the liquid mixture using a pressurized filter. Following this separation, the GVL that partitions into the aqueous phase is removed using the toluene extractor. To accurately model a liquid–liquid equilibrium (LLE) unit operation for this task, the non-random two-liquid (NRTL) model to calculate the liquid activity coefficients is modified based on LLE experiments (see the ESI† for more details). After GVL is separated from the aqueous phase, toluene and GVL are separated by distillation with GVL as a bottom product. GVL is then recycled to the biomass deconstruction section, and the toluene is reused.

3. Process synthesis

3.1. Basic design

Based on the technologies outlined in the previous section, we developed a strategy for the conversion of lignocellulosic biomass to ethanol. The process consists of ten main sections, as shown in Fig. 2: (1) feedstock handling, (2) two-stage biomass conversion, (3) liquid–liquid separation, (4) solid recovery, (5) extraction and toluene separation, (6) fermentation, (7) ethanol purification, (8) wastewater treatment, (9) GVL production, and (10) heat and power generation.

Fig. 2 Process flow diagram for the integrated biochemical ethanol production strategy. The major updates on unit operations in the work of Han et al. (2015) are highlighted by the light red boxes.

After physical pretreatment and subsequent chemical hydrolysis, a mixture of biomass deconstruction products (sugars, furfural, LA, FA, HMF, and humins) and SA in GVL–water–toluene solution (stream 8) is sent to a liquid–liquid separator (S-2) after being cooled to 298 K, where it is split into two streams: an organic (stream 10) and an aqueous (stream 13) stream; 64.4% of GVL, 26.7% of furfural, 19.2% of LA/FA, and 41.3% of HMF are recovered in the organic stream (stream 10), whereas 94.3% of sugars and 88.8% of precipitated lignin and humins remain in the aqueous stream containing the SA (stream 13), which is sent to the solid recovery section. Note that the lignin and humins remain in the solid phase in the aqueous stream (stream 13) since a large portion of GVL is recovered in the organic stream (stream 10) along with toluene.⁴⁶

Following the liquid–liquid separation, 86.3% of the organic stream (stream 11) is diverted to be recycled directly to the second stage reactor in the two-stage reaction section to maintain the required solvent mass ratio (GVL : water : toluene = 3 : 1 : 1), while stream 12 is sent to the distillation column (D-1) to separate the soluble biomass degradation products from toluene.

To avoid possible fouling, the precipitated biomass residues (humins and lignin) are removed from the solid–liquid separator (S-3) prior to being sent to the subsequent separation section and are burned in the heat and power generation section along with non-recovered GVL (1.5 wt%) and sugars (8.4 wt%) to provide heat and electricity (stream 14). The liquid mixture (stream 15) is then sent to extractor E-1 to recover the GVL remaining in the aqueous phase so that it can be reused in the conversion section. The aqueous phase containing sugars (stream 19) is sent to the fermentation section, whereas the organic phase (stream 16) is sent to the distillation column (D-1) to recover and recycle toluene.

To make up for the loss of GVL, 42.3% of the GVL-rich stream 29 from the distillation column (D-1) is sent to the catalytic bed where LA is converted into GVL using a RuSn₄/C catalyst at 493 K and 35 atm with a residence time of 5 min (R-4).^46,47 The H₂ requirement for the reaction is minimized by utilizing the internal H₂ produced from the FA decomposition. The remaining 57.7% of the bottom product in the distillation column (stream 30) bypasses the LA upgrading reactor to meet the desired GVL : water (4 : 1) ratio in the first stage reactor of the conversion section.

Prior to sugar fermentation, the stream containing sugars and SA (stream 19) is neutralized using Ca(OH)₂, and the precipitated gypsum is removed by filtration (S-4), as described in the NREL report.³⁶ C₅ and C₆ sugars in the neutralized stream (stream 20) are then diluted with a water-to-feed mass ratio of 0.2 and co-fermented to 7.9 wt% of ethanol (stream 22) over engineered yeast strains (Saccharomyces cerevisiae PE2 (stream 39)) in the fermentor (R-3) in 87% molar yields.^20,48 The ethanol (stream 22) is purified using distillation (D-2) with ethanol as a top product followed by molecular sieve adsorption (S-5) for further dehydration up to 99.5 wt%. 84.2% of the bottom liquid (stream 27) from the distillation column (D-2) is recycled to the conversion section, and the remaining fraction (stream 26) is sent to the fermentation section after wastewater treatment.

3.2. Alternative configuration

In addition to biomass residues, stream 14 contains a significant amount of sugars and GVL, so sending it directly to the heat and power generation section leads to a lower fuel yield. Accordingly, we modified the original process (solid recovery section (4) in Fig. 2) by adding water-wash (S-6) and evaporation (S-7) units after the solid–liquid separator (S-3) to recover the sugars and GVL, as shown in Fig. 3. In this configuration, the water evaporator bottom products containing water-dissolved sugars and GVL (stream 37) are sent to the extractor (E-1) along with the aqueous stream from the solid–liquid separator (stream 15). This approach increases the water loading of the downstream processing (and therefore the cost of separation) but also increases the biomass-to-fuel yield. We investigated the performance of the alternative configuration for different evaporation rates, which refer to the percentage of water removed as vapor, and found that a 95 wt% evaporation rate is optimal.

Fig. 3 Alternative process configuration for the integrated biochemical ethanol production strategy; water-dissolved sugars and GVL are supplied to the liquid–liquid extraction section.

3.3. Heat integration and power generation

A potential drawback of the proposed strategy is the high solvent : biomass mass ratio (about 11 : 1 overall) in the conversion section leading to a large heating requirement for the separation and recycling of GVL, water, and toluene. In particular, when 2000 tons of corn stover are processed per day, the total heating requirement is 171.9 MW. The energy content of the biomass residues is estimated to be 187.6 MW, which can potentially produce 133.2 MW of heat assuming that the efficiency of the combustor is 71%, which means that the total heating requirements cannot be satisfied by burning the biomass residues.

To overcome this shortcoming, we performed heat integration, including integration between process streams and unit operations, for all process configurations. For the alternative design, we were able to achieve significant energy recovery (97.1 MW), thereby reducing the heating (cooling) requirements by more than 50% to 74.8 MW (48.7 MW) as shown in Fig. 4 (see the ESI† for the major heat exchanger network and overall energy flow diagram). It is noted that the heating requirements for the reboiler of the toluene/GVL distillation column (53.1 MW) and the second-stage biomass deconstruction reactor (20.9 MW) were not fully covered by heat integration due to their high operating temperatures (430–479 K).

Fig. 4 Energy requirement before and after heat integration (alternative configuration).

The remaining heating requirements are satisfied by the heat generated from the combustion of the biomass residues. In addition to thermal fluid heating, hot combustion gases are used to produce superheated steam (783 K, 85 atm) for electricity generation. The vapor fraction of the turbine outlet is condensed to avoid compression. The turbine outlet pressure is optimized to maximize heat recovery, i.e. the outlet pressure is 1 atm in the basic design and 1.8 atm in the alternative design, so that the heating requirement for water evaporation (see section 3.2) is met by the condensation of the turbine outlet (see Fig. B.3 in the ESI†). The overall efficiency of power generation (the ratio of net power generation to biomass energy content used for power generation) is 25% for the alternative design. The surplus electricity is sold to the grid.

3.4. Assumptions and process simulation

We developed a process model using Aspen Plus, using thermodynamic parameters and reaction yields based on experimental data. To illustrate the advantages of the proposed co-solvent technology and integrated process, we perform a comparison with the GVL/water-based nonenzymatic ethanol production⁴⁰ and the enzymatic/biochemical ethanol production³⁵ strategies. Thus, we use the same assumptions and economic parameters (see Table D.1 in the ESI†), as well as the same feedstock and processing rate (2000 dry tons of corn stover per day) assuming that the reaction yields obtained herein hold for the biomass composition used in the NREL report.³⁵ However, we note that adopting the assumptions and parameters from more recent reports^34,49 would lead to higher capital costs for all strategies (see Table D.2 in the ESI†).

The equipment cost of the liquid–liquid separation, extraction/toluene separation, GVL production, and a part of the two-stage reaction (including pumps) sections was estimated using the Aspen Process Economic Analyzer.⁵⁰ The equipment cost of the remaining sections was estimated using a scaling expression based on the equipment size and cost data in the NREL report.³⁵ The equipment and utility costs of the heat exchanger network were calculated using the Aspen Energy Analyzer. All equipment and material costs were adjusted to a common basis year of 2007.⁵¹ The minimum selling price (MSP), i.e. the ethanol selling price that makes the net present value zero, was calculated using the discounted cash flow methodology used by the NREL³⁵ (see details in the ESI†). In the succeeding section, we will refer to the proposed GVL/water/toluene-based nonenzymatic ethanol production, the GVL/water-based nonenzymatic ethanol production,⁴⁰ and the enzymatic/biochemical ethanol production³⁵ strategies as strategy A, strategy B, and strategy C, respectively.

4. Results

4.1. Capital and operating costs

Table 2 shows the capital cost breakdown of the basic and alternative designs of strategy A and a comparison with strategies B and C. The total installed equipment cost is computed to be $202.1 million for the alternative design, while the total project investment, which includes other direct (e.g. warehouse and site development) and indirect (e.g. field expenses, home office and permit costs, and project contingency) costs, is estimated to be $371.9 million. The fuel production rates for the different strategies are summarized in the ESI.†

Table 2 Capital cost components ($ per GGE) for different strategies

Strategy	Strategy A		Strategy B (Han et al., 2015)	Strategy C (Humbird et al., 2011)
	Basic	Alternative
Pretreatment/biomass conversion	0.267	0.250	0.240	0.107
Extraction/toluene separation	0.025	0.023
Extraction/CO2 separation			0.160
Biochemical conversion	0.021	0.021	0.021	0.161
Catalytic conversion	0.002	0.001	0.001
Distillation/solid recovery	0.059	0.065	0.063	0.073
Wastewater treatment	0.116	0.112	0.117	0.161
Heat and power generation	0.226	0.199	0.189	0.215
Utilities	0.015	0.013	0.016	0.022
Storage	0.016	0.016	0.016	0.016
Total annualized installed equipment cost per GGE	0.748	0.700	0.823	0.756

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The total annualized installed equipment cost of the alternative design is lower than that of the basic design mainly because the alternative configuration (1) has a higher fuel production rate (35.7 × 10⁶ GGE per year vs. 38.4 × 10⁶ GGE per year) and (2) requires a smaller heat and power generation section (204.8 MW vs. 187.6 MW), although the cost for distillation and solid recovery is higher due to the additional water-wash and evaporation steps for sugar and GVL recovery. Compared to strategy B, the total annualized installed equipment cost ($ per GGE) of the alternative design is 14.9% lower because the new separation steps are less expensive than the CO₂-based separation ($6.8 million vs. $47.8 million), which consist of multiple compressors and vapor–liquid separators operating at 67 atm, despite a 3.5% lower fuel production rate. The capital cost for the alternative design is also lower than that of strategy C primarily because it (1) does not require on-site cellulase enzyme production or enzymatic hydrolysis steps, (2) has smaller wastewater treatment resulting from lower stillage and condensate production during biomass deconstruction and ethanol purification (411 178 kg h⁻¹vs. 202 034 kg h⁻¹ of wastewater), and (3) has a smaller heat and power generation section resulting from smaller by-product streams. The capital cost is lower despite the fact that the cost of the biomass pretreatment in the alternative design is higher than that in strategy C and the fuel production rate of the alternative design is slightly lower than that of strategy C (38.4 × 10⁶ GGE per year vs. 40.2 × 10⁶ GGE per year).

Table 3 shows a comparison of the operating costs for the four strategies. The feedstock cost has the largest contribution in all the strategies. The total operating cost of the alternative design is 7.1% lower than that of strategy C mainly due to the absence of enzyme production ($0.294 per GGE) and ammonia for hydrolysate conditioning (included in other raw materials, $0.109 per GGE), despite a 4.4% lower fuel production rate, higher waste disposal cost resulting from a large amount of gypsum precipitating in liming ($5.3 million per year vs. $1.5 million per year), and additional material cost for toluene ($0.023 per GGE) and RuSn₄/C (included in other raw materials, $0.010 per GGE). The total operating cost of the alternative design is also slightly lower than that of strategy B primarily because of the lower fixed costs (e.g. salaries, maintenance and property insurance costs), which are calculated based on the total installed equipment cost. The large difference in total operating costs of the basic and alternative designs is mainly due to the difference in the fuel production rates (about 7.0% lower in the basic design).

Table 3 Operating cost components ($ per GGE) for different strategies

Strategy	Strategy A		Strategy B (Han et al., 2015)	Strategy C (Humbird et al., 2011)
	Basic	Alternative
a Waste disposal costs include gypsum disposal following sulfuric acid neutralization during fermentation and other solid waste disposal (e.g. ash removal from the boiler).
Glucose				0.294
Toluene makeup	0.016	0.015
CO2 makeup			0.033
Sulfuric acid makeup	0.074	0.069	0.068	0.037
Other raw materials	0.235	0.224	0.226	0.296
Waste disposal^a	0.155	0.137	0.131	0.037
Catalyst regeneration	0.010	0.007	0.005
Utility usage	0.026	0.012	0.015
Fixed costs	0.284	0.268	0.304	0.266
Feedstock	1.265	1.178	1.136	1.126
Total operating cost	2.066	1.910	1.918	2.057

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4.2. Minimum selling price

Fig. 5 compares the MSPs, calculated using discounted cash flow analysis, where the additional revenue obtained through the sale of surplus electricity is represented by a negative sign. The lowest MSP is $2.95 per GGE for the alternative design of strategy A. Compared to the MSP of the basic design ($3.11 per GGE), the MSP of the alternative design is improved due to a higher fuel production rate, thus outweighing the capital and operating costs of the additional operations. We note that the strategies reported herein (in particular, the alternative design) result in better MSP values than strategy B ($3.37 per GGE). This improvement is because the average total annualized cost of the alternative design ($122.8 million per year) is 10.5% lower than that of strategy B ($137.2 million per year) due to the less expensive toluene-based solvent/solid recovery system operating at atmospheric pressure. Moreover, the revenue obtained from excess electricity sales ($0.25 per GGE) is higher than that of strategy B ($0.08 per GGE). In addition, the MSP for the alternative design is also better than that of strategy C because of the 11.5% lower average total annualized cost resulting from the absence of expensive enzymatic hydrolysis and on-site enzyme production steps and 51.8% higher net electricity surplus, despite 4.4% lower ethanol production.

Fig. 5 Comparison of costs and revenues of the different strategies ($ per GGE). MSP: minimum selling price.

Fig. 6 shows the contribution of individual sections toward the MSP on a $ per GGE basis. The feedstock cost is the highest contributor (included in the feedstock handling section), while the two-stage reaction section is the second highest contributor. The contribution of the capital cost to the reaction cost is found to be significantly higher than those of material and other operating costs.

Fig. 6 Cost contribution per process section ($ per GGE).

4.3. Further discussion

Since the heating requirements of the proposed strategy are significant, the credit from surplus electricity sales is low compared to the total production cost while the cost of the turbo generator is substantial (12.1% of the total installed equipment cost). To study this trade-off, we design an alternative configuration, in which there is no turbo generator and the electricity requirement is met externally and the excess biomass residues are disposed. Fig. 7 shows how the MSPs of the two strategies change with electricity selling price. If the electricity selling price is higher than $0.05 kWh⁻¹ (region A), the installation of a turbo generator is better than the disposal of biomass residues, while the disposal of biomass is a better option if the electricity price is less than $0.05 kWh⁻¹ (region B). It is important to note that although in region B electricity generation using residual biomass is an inferior option from an economical point of view, it leads to lower greenhouse gas (GHG) emissions. We also note that the economics of the option to use a turbo generator can be improved if the renewable identification number (RIN) credit is considered.

Fig. 7 Minimum selling price of two different residual biomass handling strategies. Electricity generation is a better option in region A, while disposal of excess biomass residues is better in region B.

We also studied the impact of changes in the discount rate and tax rate (Fig. 8) using the values reported recently in the literature, i.e. we decreased the discount rate from 10% to 6.74% and the tax rate from 35% to 13%.^52–54 The discount rate has a significant effect on the MSP, while the effect of the tax rate is moderate, because the former affects the annualized capital cost, which is one of the dominant cost drivers. Finally, when all parameters were adjusted simultaneously, the MSP for the proposed strategy decreases to $2.57 per GGE.

Fig. 8 Impact of adjusting the key economic parameters on minimum selling price ($ per GGE).

5. Conclusions

We developed a novel strategy for the production of ethanol from lignocellulosic biomass by integrating a nonenzymatic chemical hydrolysis technology (using a mixture of GVL/water/toluene as a solvent containing a dilute sulfuric acid catalyst) with new separation subsystems. In addition to the basic design, we proposed an alternative configuration that leads to a higher ethanol production yield. It was shown that the process heating and electricity requirements can be satisfied by the combustion of the residual biomass after heat integration. The combined energy available from the ethanol product and excess electricity is 44.1% of the energy content of the biomass feed. The technoeconomic analysis showed that the MSP of the proposed strategy can be reduced to $2.95 per GGE, which is 12.5% lower than that achieved by the GVL/water-based nonenzymatic ethanol production⁴⁰ (strategy B), primarily due to the replacement of the expensive CO₂-based sugar recovery section and the energy-intensive GVL evaporator with solid recovery with simple and less expensive toluene-based extraction and filtration processes. The MSP of the proposed strategy is also 9.7% lower than that of the benchmark enzymatic strategy³⁵ (strategy C). Finally, we note that the proposed sugar production process can be integrated with other conversion technologies for the production of different bioproducts such as butanol, gasoline and diesel range fuel blendstocks.

Acknowledgements

This material is based on the work supported by the DOE Great Lakes Bioenergy Research Center (DOE Office of Science BER DE-FC02-07ER64494).

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Footnote

† Electronic supplementary information (ESI) available. See DOI: 10.1039/c6re00227g

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