The Techno-Economic Basis for Coproduct Manufacturing To Enable

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The techno-economic basis for coproduct manufacturing to enable hydrocarbon fuel production from lignocellulosic biomass Mary J. Biddy, Ryan Davis, David Humbird, Ling Tao, Nancy Dowe, Michael T. Guarnieri, Jeffrey G Linger, Eric M. Karp, Davinia Salvachua , Derek Richard Vardon, and Gregg T Beckham ACS Sustainable Chem. Eng., Just Accepted Manuscript • DOI: 10.1021/ acssuschemeng.6b00243 • Publication Date (Web): 27 Apr 2016 Downloaded from http://pubs.acs.org on May 4, 2016

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ACS Sustainable Chemistry & Engineering

The techno-economic basis for coproduct manufacturing to enable hydrocarbon fuel production from lignocellulosic biomass Mary J. Biddy*, Ryan Davis, David Humbird, Ling Tao, Nancy Dowe, Michael T. Guarnieri, Jeffrey G. Linger, Eric M. Karp, Davinia Salvachúa, Derek R. Vardon, Gregg T. Beckham* National Bioenergy Center, National Renewable Energy Laboratory, 15013 Denver West Parkway, Golden CO 80401, USA * Corresponding authors: [email protected]; [email protected] Abstract: Biorefinery process development relies on techno-economic analysis (TEA) to identify primary cost drivers, prioritize research directions, and mitigate technical risk for scale-up through development of detailed process designs. Here, we conduct TEA of a model 2,000 dry metric ton-per-day, lignocellulosic biorefinery that employs a two-step pretreatment and enzymatic hydrolysis to produce biomass-derived sugars, followed by biological lipid production, lipid recovery, and catalytic hydrotreating to produce renewable diesel blendstock (RDB). Based on projected, near-term technical feasibility of these steps, we predict that RDB could be produced at a minimum fuel selling price (MFSP) of USD$9.55/gasoline-gallon-equivalent (GGE), predicated on the need for improvements in the lipid productivity and yield beyond current benchmark performance. This cost is significant given the limitations in scale and high costs for aerobic cultivation of oleaginous microbes and subsequent lipid extraction/recovery. In light of this predicted cost, we develop an alternative pathway which demonstrates that RDB costs could be substantially reduced in the near term if upgradeable fractions of biomass, in this case hemicellulose-derived sugars, are diverted to coproducts of sufficient value and market size; here, we use succinic acid as an example coproduct. The coproduction model predicts an MFSP of USD$5.28/GGE when leaving conversion and yield parameters unchanged for the fuel production pathway, leading to a change in biorefinery RDB capacity from 24 MM GGE/year to 15 MM GGE/year and 0.13 MM tons of succinic acid per year. Additional analysis demonstrates that beyond the near-term projections assumed in the models here, further reductions in the MFSP towards $2-3/GGE (which would be competitive with fossil-based hydrocarbon fuels) are possible with additional, transformational improvements in the fuel and coproduct trains, especially in terms of carbon efficiency to both fuels and coproducts, recovery and purification of fuels and coproducts, and coproduct selection and price. Overall, this analysis documents potential economics for both a hydrocarbon fuel and bioproduct process pathway, and highlights prioritized research directions beyond the current benchmark to enable hydrocarbon fuel production via an oleaginous microbial platform with simultaneous coproduct manufacturing from lignocellulosic biomass. Keywords: Biochemicals | lignocellulose | integration | biorefinery | Biofuels INTRODUCTION Lignocellulosic biomass offers significant promise to offset a large portion of mankind’s fossil fuel usage as part of a balanced renewable energy portfolio for transportation fuels, renewable chemicals, robust materials, and as a partial source of sustainable power generation. Among all major renewable energy technologies examined to date, biomass is the primary source of 1-2 renewable liquid fuels for vehicle, air, and maritime transportation. As highlighted in a number of studies, including the SCOPE 3 Bioenergy and Sustainability study, well over 50 countries have biofuels mandates or blending targets. These mandates are 3-4 dominated by ethanol, with typically smaller biodiesel mandates, in the range between 1-20% of the transportation pool. The vast majority of ethanol production today comes from starch or sugar cane-based industries, which exhibit a substantial scale 5 limitation relative to worldwide fuel demand and has caused significant concerns regarding offsetting food production for fuels. As a means to ameliorate the concerns regarding the food versus fuel debate, further reduce greenhouse gas emissions, support energy security, promote rural job development, and to achieve fuel scale production beyond starch- and sugarcane based feedstocks, ethanol production from lignocellulosic biomass is under intense development at the industrial scale. The current predominant route for lignocellulosic ethanol production proceeds via biochemical conversion processes that conventionally includes a mild thermochemical pretreatment or fractionation step, enzymatic hydrolysis with cellulolytic enzymes, and 6-7 fermentation by an ethanologen such as Saccharomyces cerevisiae or Zymomonas mobilis. Certainly many variations exist on 8-9 that theme including, for example, many different types of pretreatment and process configurations with separate or 10-11 combined hydrolysis and fermentation. Multiple techno-economic analysis (TEA) and life-cycle analysis (LCA) studies have been conducted for various configurations of industrial cellulosic ethanol plants, all of which point towards a path to produce ethanol that is at, or close to, being competitive with petroleum derived fuels with the potential ability to offset substantial 12-25 greenhouse gas (GHG) emissions, and indeed several cellulosic ethanol plants are currently being brought online worldwide. Going forward, however, ethanol is a renewable additive primarily for gasoline-range fuels, but it is not directly applicable to a broader pool of transportation fuels including diesel, jet, and maritime fuels, nor does it typically suffice as a complete substitute for gasoline in most engines, a limit typically referred to as the “blend wall”. Lastly, ethanol distribution does not follow the same channels as petroleum-derived fuels. The issues of the limited fuel market capacity, the need for another fuel distribution system, and the ethanol blend wall have thus motivated research efforts to move towards the production of “drop-in”, infrastructure compatible hydrocarbon fuels.

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Drop-in, lignocellulosic biofuels beyond ethanol that are derived from biological or catalytic routes have been examined primarily at the laboratory or pilot-scale, with many in scientific literature first focusing on the development of processes starting from “clean” sugars. Many challenges remain in both biological and catalytic routes to hydrocarbon fuels, primarily related to realizing integrated processes that are able to completely convert lignocellulosic biomass into a finished fuel. For example, 35-37 inorganic components found in biomass, which are typically difficult to remove, oftentimes poison chemical catalysts. Biological production of hydrocarbons from biomass-derived sugars typically suffers from the challenges of insufficient titers, productivities, and yields when a pathway is first introduced as well as toxicity effects stemming from inhibitors produced during 38-40 biomass deconstruction. These challenges, among many others in process integration and yield improvements, must be overcome to cost effectively produce hydrocarbon biofuels from lignocellulosic biomass. To bring hydrocarbon fuel production to a level of technological viability similar to that of ethanol production, TEA and LCA studies will be key to consistently evaluate and compare the myriad process options that exist and are being constantly proposed and developed. To that end, Davis et al. recently published a peer-reviewed, publically available report describing a projected TEA and LCA model that examines the cost of producing hydrocarbon biofuels via a biological upgrading step from lignocellulosic 41 sugars and sustainability considerations within the bioefinery conversion facility. The primary purpose of this model was to identify research gaps and cost drivers to achieve a near-term 2017 technical target to produce hydrocarbon fuels at a cost of USD$5 per gasoline gallon equivalent ($5/GGE). The modeled biochemical conversion process described in that work employs a 42 2-step pretreatment process consisting of a deacetylation step followed by dilute-acid pretreatment , enzymatic hydrolysis with cellulase enzymes, and an aerobic cultivation step to employ an oleaginous microbe to produce free fatty acids (FFAs), which are 41 secreted into the broth, recovered, and catalytically upgraded to a renewable diesel blendstock (RDB). The predicted RDB cost from this process is $5.10/GGE in 2011-USD dollars (hereafter the basis for cost figures unless otherwise noted). The main cost drivers, and thus identified research needs, associated with realizing a process based on this model include achieving FFA yield and productivity of 0.28 g/g sugars and 1.3 g/L/hr, respectively coupled to the ability to (1) secrete FFAs and (2) convert a mixed sugar stream to FFAs. The productivity target was chosen at the time for consistency with what is achievable in ethanol 22 41 fermentation and the yield is within reach from literature values for FFA yields from glucose. However, we note that both the previous productivity assumption and FFA secretion have yet to be achieved to our knowledge in an oleaginous microbe. FFA 43 secretion has been achieved in several non-oleaginous microbes including the cyanobacteria Synechocystis sp. PCC6803 , 44 45-47 Escherichia coli , and Saccharomyces cerevisiae , suggesting that fatty acid secretion is a distinct possibility in oleaginous microbes. Moreover, the scale of reactors considered previously did not account for the substantial differences in aerobic and anaerobic bioreactor cultivation of microbes. As will be demonstrated here, these assumptions have major effects on the estimated cost of an oleaginous fuel platform. The cost projection of $5/GGE would represent a milestone towards the ultimate objective of producing hydrocarbon fuels at 41 prices concomitant with conventional petroleum fuels today (e.g., ~$2-3/GGE). To achieve a lower price, significant process modifications will likely need to be considered. One option to significantly lower the cost of fuel production from biomass, as is practiced throughout the petroleum refining industry, is to rely on the production of high-value coproducts alongside fuels, as 34, 48-59 has been noted throughout the biomass valorization literature. Production of petrochemicals in many cases employs oxidation of hydrocarbon platform molecules, and, relative to petroleum, biomass-derived platform molecules from polysaccharides and lignin already exhibit significant oxygen content, making them, from a stoichiometric perspective, more 53, 59 attractive building blocks for oxygen-containing chemicals if appropriate molecular transformations can be realized. The production of bio-based chemicals, however, also presents a substantial set of technical challenges. For example, significant emphasis has been placed in pioneering reports on the production of single platform molecules, often from sugars for uses such 48, 50, 53 as chemical and polymer precursors. Especially for polymer and specialty chemical applications, precursor purity is often of utmost importance, which in an integrated process from lignocellulosic biomass will require scalable, cost-effective separations and cleanup strategies. As a result, it has also been suggested that bio-based chemicals should rely on the intrinsic heterogeneity and intrinsic polymeric structure already present in biomass to make functional replacement materials rather than 56 depolymerization to monomers that can then go on to be converted to building blocks. Regardless, whether biomass-derived substrates are converted to new platform molecules or used directly as polymers, process integration issues including separations, cleanup, and quality of the final products will be of paramount importance to replace products used by consumers currently derived from petroleum. Moreover, the market size disparity between fuels and chemicals will be a key driver in the attainable scale for a single process, further warranting the need to develop modular processes that can be adjusted to produce several different chemicals, either spread across multiple biorefineries, or within the same biorefinery, depending on the economies of scale for a given process configuration or coproduct. Overall, for the production of hydrocarbon biofuels from lignocellulose, the basis for manufacturing coproducts alongside fuels must be quantified with TEA to ascertain the relative contributions that coproducts can make to the overall economic viability. To that end, here we present TEA models for several process scenarios centering on the core concept of a biorefinery wherein hydrocarbon fuels and chemicals are produced in parallel from lignocellulosic biomass, similar to the simultaneous production of


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fuels and chemicals practiced in petroleum refining. The baseline process considers economics of fuel production alone, similar to that noted above as presented in Davis et al. and is a 2-step pretreatment approach, enzymatic hydrolysis, biological 41 production of lipids, and a catalytic step to produce RDB from both hemicellulose and cellulose-derived sugars. However, for this case a number of updates are made to reflect more realistic experimental benchmarks (such as intracellular lipid productivity metrics via oleaginous yeasts rather than more idealized high-productivity FFA-secretion), and to incorporate improved design details (such as the use of revised design/cost estimates for bubble column bioreactors). We then examine a coproduct manufacturing scenario wherein succinic acid is produced predominantly from hemicellulose-derived sugars fractionated after dilute-acid pretreatment and the cellulose-derived sugars are used to produce lipids for RDB. Succinic acid is chosen as a representative molecule to demonstrate the potential for coproduct pathways within this framework for an integrated biorefinery, with preliminary research demonstrating biological production of succinic acid at encouraging performance levels. When viewed as an intermediate molecule, there is a strong potential for a greater demand of succinic acid given that there are numerous known technologies for the production of a wide slate of high value, large market commodity chemicals from succinic 48, 53, 60 acid. The model predicts that the production price of hydrocarbon fuels can be significantly improved by the simultaneous production of chemicals from other biomass fractions. Key findings from this work regarding process selection are: (1) maximizing carbon efficiency, especially to higher-value coproducts is paramount, (2) recovery and purification of both fuels and coproducts is a major cost driver that will benefit significantly from research and development activities to improve applicability at the biorefinery scale, and (3) the product selection and price will greatly influence the process flexibility and the ability to balance process severity. Overall, this study quantifies the economic baseline for the manufacturing of next-generation, hydrocarbon fuel biorefineries, demonstrates that co-products can enable the economics of hydrocarbon fuel production, and provides a framework for further research and development in a next-generation integrated lignocellulosic biorefinery. SUMMARY OF MODEL As illustrated in Figure 1A, the base case design produces hydrolysate via deacetylation and dilute-acid pretreatment of biomass followed by enzymatic hydrolysis. Unconverted biomass solids (mostly lignin) are removed from the hydrolysate through a flocculant-aided solid/liquid separation, and the remaining liquor fraction is concentrated prior to biological upgrading. The hydrolysate, which is rich in both C5 and C6 sugars, is biologically converted in an aerobic process to lipids that are stored intracellularly by an oleaginous microbe. The lipid fraction is extracted from the cells in a wet extraction process and subsequently upgraded in a hydrotreating unit to produce RDB. Supporting processes, including wastewater treatment, utilities (including cooling water, chilled water plant and instrument air, and process water), and the combined heat and power system (consisting combustor, boiler, and turbogenerator) are also incorporated in the design and economics. A variation of the process design, as outlined in Figure 1B, considers the conversion of a separate C5-rich hydrolysate stream to a chemical coproduct. In this case, an additional solid/liquid separation step is introduced after the pretreatment process and the C5-rich liquor is sent to a biological upgrading process for continuous conversion to succinic acid, followed by the required separation and purification steps. The solids (C6-rich hydrolysate stream) continue to enzymatic hydrolysis and follow the same process steps to RDB as outlined in the base design. Additional details of the design and the basis for both the “current experimental benchmark” and an “improved” case are summarized in the following sub-sections. The base case process (whole hydrolysate to fuels alone) is discussed first, followed by the modifications made for the alternative coproduction of chemicals and fuels process (C6-to-fuels, C5-to-coproducts). Pretreatment. Milled biomass, with a composition described in Table 1, enters the process facility and is assumed to be “throatof-reactor” ready, i.e. any pre-processing required prior to the first unit operation is included in the delivered feedstock cost. The biomass is conveyed to the pretreatment area where it is first processed in a deacetylation step that uses a sodium hydroxide soaking process (loaded at 17 mg of NaOH/g of dry biomass) at 80°C for 1 hour in a set of parallel batch reactors. This process step has been shown to remove various components of the biomass that are either potentially toxic in downstream biological upgrading or inert biomass components, including acetate, lignin, extractives and ash, and improves overall yields and reduces 41-42, 61 The solubilized fraction of the biomass is drained from the batch reactors capital cost due to lower process throughputs. and sent to wastewater treatment. Overall, this liquor fraction consists of a range of biomass components with solubilization and removal of 100 wt% of the extractives, 75 wt% of the ash, 20 wt% of the lignin, 2 wt% of the xylan, 50 wt% of the sucrose, and 88 wt% of the acetate. The conversion in the deacetylation step is consistent for all of the cases studied. The remaining solids are conveyed to the dilute acid pretreatment reactor at a total solids loading of 30 wt%, where a large fraction of the hemicellulose carbohydrates are converted to soluble sugars including xylose, arabinose, galactose, mannose and 41 some glucose. The basis for the design of the reactor system is described in a previous study. The biomass is treated with 9 mg H2SO4/g dry biomass (based on acid loading present in the pretreatment reactor), with a total residence time in the reactor of 5 minutes at a temperature of 165°C. The overall conversion reactions are provided in Table 2 and summarize both the current 61 benchmark case as demonstrated in NREL pilot plant runs and projected improved case. Following the acid pretreatment, the


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hydrolysate whole slurry containing 30 wt% total solids and 16 wt% insoluble solids is neutralized with ammonia in stoichiometric quantities to raise the hydrolysate pH to 5 and the stream is diluted with water to a total solids concentration of 20 wt% prior to enzymatic hydrolysis.

Figure 1: Overview of design block flow diagrams from (A) baseline fuels only design and (B) coproduction of fuels and chemicals design. Table 1: Delivered Feedstock Composition Assumed in the Process Design Component Glucan Xylan Lignin Ash Acetatea Protein Extractives Arabinan Galactan Mannan Sucrose Total structural carbohydrate Total structural carbohydrate + sucrose Moisture (bulk wt %) a

Composition (dry wt %) 35.05 19.53 15.76 4.93 1.81 3.10 14.65 2.38 1.43 0.60 0.77 58.99 59.76 20.0

Represents acetate groups present in the hemicellulose polymer; converted to acetic acid in pretreatment.


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ACS Sustainable Chemistry & Engineering Table 2: Overall Yields for Benchmark and Improved Case (Biomass Deconstruction and Fuel Production Operations)

Experimental Benchmark Case

Improved Case

Pretreatment Solids loading (%) 30% 30% Acid loading (mg H2SO4/g dry biomass present in reactor) 9 9 10% 10% % Conversion: (Glucan)n + n H2O→ n Glucose % Conversion: (Xylan)n + n H2O→ n Xylose 76% 78% 90 % 90% % Conversion: (Galactan)n + n H2O→ n Galactose % Conversion: (Arabinan)n + n H2O→ n Arabinose 90% 90% % Conversion: Acetate → Acetic Acid 100% 100% % Conversion: Lignin → Soluble Lignin 5% 5% S/L Separation: Soluble sugar loss to solids 4% 1% Enzymatic Hydrolysis + Hydrolysate Conditioning Enzyme loading (mg/g cellulose) 12 10 Enzymatic hydrolysis time (days) 5.0 3.5 % Conversion: (Glucan)n + n H2O → n Glucose 86% 90% 93% 93% % Conversion: (Xylan)n + n H2O→ n Xylose S/L Separation: Soluble sugar loss to solids 5% 5% Bioconversion (Lipid Production) Lipid productivity (g/L-hr) 0.34 0.4 Lipid content (wt%) 60% 70% % Conversion: Glucose → Lipid [total glucose utilization] 1 75% [100%] 82% [100%] % Conversion: Xylose → Lipid [total xylose utilization] 1 44% [59%] 80% [98%] Bioconversion metabolic yield [process yield] (g/g) 2 0.25 [0.24] 0.27 [0.27] Lipid Recovery + Upgrading Extraction pretreatment acid loading (wt% of feed liquor) 1% 1% Extraction solvent loading (g hexane/g cell mass) 5 5 Extraction net lipid recovery 90% 90% Hydrotreating fuel yield (g RDB + naphtha/g lipid) 0.81 0.81 Hydrotreating net H2 makeup demand (wt% lipid feed) 1.8% 1.8% 1 First number represents sugar conversion to desired product (lipids), values in brackets indicate total sugar utilization (including diversions to cell mass) 2 Metabolic yield = g product per g sugars converted; process yield = g product per g total sugars available for bioconversion

Enzymatic hydrolysis. In the base case, the whole slurry hydrolysate stream is fed at a 20 wt% total solids to a high-solids continuous reactor for enzymatic hydrolysis by a cellulase enzyme prepared on-site. The partially hydrolyzed slurry is next batched to one of several parallel bioreactors. Hydrolysis is completed in the batch reactor with an overall hydrolysis time of 3.5 days in the continuous and batch steps. This process converts both glucan and residual xylan in the solids to monomeric glucose and xylose, respectively. Table 2 summarizes the key process conditions and yields for both the benchmark and improved cases for the enzymatic hydrolysis process. Hydrolysate conditioning: S/L separations and concentration. The slurry is further conditioned for downstream biological upgrading by removal of insoluble solids required for effective aerobic biological conversion and subsequent product purification. A vacuum filter press is used for this solid/liquid separation step. To facilitate and improve filtration, polymer flocculants at a loading of 10 g/kg insoluble solids are added to reduce the overall required filter area. The process conditions for both the benchmark and improved cases are outlined in Table 2 for this step, with the benchmark values based on data 62 demonstrated from bench scale studies reviewed in Sievers et al. The solids removed in the filter press are sent to the combustor/steam turbogenerator, which provides all process heat and a portion of the power requirements for the integrated process design. The clarified soluble sugar stream is split with roughly 15% routed directly to bioconversion to initiate fed-batch fermentation, and the majority (85%) concentrated from 20 to 50 wt% total sugars in a mechanical vapor recompression (MVR) evaporation system. The concentrated sugar stream exiting the evaporators is cooled and fed to the bioreactors. Biological Conversion. The biological conversion step consists of a seed train of increasingly larger stirred tank reactors in series that are dedicated to biomass seed propagation (grown with a 10% split of the starting available hydrolysate sugars), and a system of aerobic bubble columns for the production of lipids. Many experimental efforts in the literature have focused on the 63 development and evaluation of intracellular lipid accumulation in model organisms such as Saccharomyces cerevisiae or 64-66 67-69 70-72 Escherichia coli or oleaginous microbes, such as Yarrowia lipolytica , Rhodosporidium toruloides , and Lipomyces 72-74 74 starkeyi , among many others. The biomass composition in the current model is based on L. starkeyi. The production


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bioreactors are modeled as bubble columns each with 500 m (500,000 L) of total internal volume operated in a fed-batch mode 3 of operation, with an initial 50% fill level up to a final 80% fill level (400 m maximum ungassed working volume) over the course of the fed-batch cycle. The reactors are cooled through an external pump-around loop circulated through a chilled water heat exchanger. Key cost and design considerations for this system are attributed to the aerobic nature of the bioconversion step, with submerged aerobic production of the yeast cells dependent on the target lipid productivity rate (g/L-hr) and the associated oxygen uptake rate (OUR), which is set in the model to be equal to the oxygen transfer rate (OTR). The required OTR is achieved in the model by varying the aeration rate to the bubble columns, given a fixed vessel geometry and air compressor pressure (liquid height plus hydraulic losses). Further details on vessel design and operation, as well as underlying drivers behind the OTR/OUR in the modeled system, are provided in the Supplemental Information. Table 2 provides a summary of key conversion and yield parameters for the bioconversion step attributed to the experimental benchmark and projected improved cases. Lipid separations. The hydrocarbon intermediate product is composed of intracellular lipids (primarily triglycerides). The lipid fraction is extracted from the cells in a wet extraction process. The extraction step follows a similar schematic as published for 75-76 extraction of lipids from algal biomass with demonstrated yields up to 92.5% , consisting of a dilute acid treatment operation targeting 1% acid loading at 150°C, followed by cooling, flashing, and neutralization (all steps being similar to the initial biomass pretreatment step discussed previously). The material is then routed to a counter-current extraction column with hexane solvent at a total 5:1 solvent loading (solvent versus yeast dry cell weight). The extract consisting of primarily solvent and lipids is sent to a stripping column to recover the hexane, while the aqueous raffinate product is sent to wastewater treatment. The key process conditions and yields for this step are summarized in Table 2. Upgrading and recovery. Finally, the recovered lipid product exiting the extraction/separation step is upgraded via hydroprocessing to a hydrocarbon product, primarily RDB with a small naphtha coproduct. This operation is designed for deoxygenation and saturation of the lipid material, yielding a paraffinic end-product suitable for blending. The process conditions 77 and yields for this step are summarized in Table 2, and are based on detailed experimental data documented in Marker et al. for the upgrading of vegetable oils. The lipid product is modeled as a representative triglyceride component consisting of oleic (18:1), stearic (18:0), and palmitic (16:0) acid chains, which constitute the majority of the fatty acid profile characterized in many 74 oleaginous yeast. Based on this representative component, plus the targeted 81 wt% fuel yield (79% RDB plus 2% naphtha versus lipid feed), the modeled hydrogen consumption is roughly 1.8 wt% of the lipid feed to the upgrading step. Final fuel blendstock yields following the upgrading section are calculated based on the total energy yield and are converted to a GGE basis 78 according to lower heating values (LHV), assuming standard published LHV values for gasoline. Off-gases from the upgrading section are combined and routed to the boiler, while produced water is sent to wastewater treatment. Modifications to the base design for the coproduction of fuels and chemicals. While the base case design is focused solely on the production of fuels in the biorefinery, an alternative scenario is investigated that considers the economic impact of the coproduction of chemicals and fuels. For this new basis with coproduction, the benchmark design is modified to separate and upgrade the C5-rich hydrolysate stream to chemicals via biological conversion, while the residual C6-rich slurry is upgraded to fuels. More specifically, the hydrolysate slurry after dilute acid pretreatment is immediately processed in a vacuum filter to separate solubilized hemicellulose sugars from cellulose and residual unconverted solids. The current benchmark case for this solid/liquid separation step is based on bench-scale demonstrations on corn stover-derived pretreated hydrolysate. It was demonstrated that recoveries of 96 wt% of the sugars using a wash ratio of 12.4 L/kg (wash water versus liquor in the filter cake) and a filter capacity of 15.8 kg insoluble solids (IS)/m2h could be achieved. The target case utilizes process assumptions based on previously 79 reported experimental results outlined in Sievers et al. In that study, a hydrolysate produced by dilute acid pretreatment, but without deacetylation, was separated with a sugar recovery efficiency of 99 wt% utilizing a wash ratio of 4.3 L/kg and filter 2 capacity of 20 kg IS/m -hr. The recovered solids that are rich in unconverted cellulose are sent to the enzymatic hydrolysis step and further upgraded to fuels, whereas the liquid, which will be referred to as the C5-rich liquor, is sent to the C5 conversion train for the production of chemicals. The solids fraction is routed to enzymatic hydrolysis at a 17.5 wt% total solids loading, which differs from the concentration commonly employed in whole-slurry processing and that was discussed earlier for the base case. Since the coproduction of chemicals and fuels pathway employs a washed-solids hydrolysis whereby a large fraction of soluble solids (namely hemicellulose sugars) removed, the overall viscosity of this stream is increased due to the higher fraction of insoluble solids relative to total solids. As the material approaches 20% IS, it becomes considerably more difficult to pump and agitate in the early stages of hydrolysis, often resulting in reduced glucose yields. The total solids loading is therefore reduced to 17.5 wt% for this modified design. However, the downstream yields and all process design assumptions for the enzymatic hydrolysis and conversion/upgrading steps are consistent with the fuel-only base case.


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The C5-rich hydrolysate liquor continues to a parallel biological upgrading train for chemical coproduction. For this design, succinic acid has been identified as an example target chemical product, utilizing a process modeled with Actinobacillus succinogenes as the production organism. This organism can utilize pentose and hexose sugars, as well as hydrolysates, with the 80-81 Succinic acid is meant to serve as a representative production of succinic acid being greatly enhanced by the presence of CO2. molecule to demonstrate the potential enabled by introducing coproduct pathways within this integrated biorefinery framework. This integrated biorefinery concept should thus be applicable to the coproduction of fuels and a wide of array of other chemical intermediates. In this design, the C5 liquor is initially split, with 10% of the liquor being sent to the seed fermentor used to grow A. succinogenes. The other 90% of the liquor is sent to the continuous biological reactor system. CO2 is also fed to the bioreactor at 82 a rate of 0.1 VVM, which is adopted from literature. Sodium hydroxide is added to the biological conversion process to maintain the pH of the system near 7, with stoichiometric amounts of NaOH added to the fermentation broth. The near-term conversion for the bioreactor for the “improved” case are set at 0.73 g succinic acid/g pentose sugars (xylose and arabinose) and 0.73 g succinic acid/g hexose sugars (glucose and galactose), with pentose sugar utilization targeted to improve to similar levels 83 as those reported publicly for hexose sugars. Selectivity to acid byproducts is assumed to be minimized to 0.078 g/g C5 sugars and 0.081 g/g C6 sugars for acetic acid coproduction and 0.02 g/g C5 sugars and 0.03 g/g C6 sugars for formic acid 83 coproduction. The succinic acid productivity rate is targeted to be 2.0 g/L-hr, which is realistic for the continuous fermentation pursued for this design. Recent results have demonstrated that succinic acid can be produced from corn stover hydrolysates at productivities of 1.45 g/L-hr utilizing a continuous fermentation process and overall metabolic yields at 0.78 g succinate/g of 84 sugar consumed. These metrics are the values incorporated for the experimental benchmark case. Since A. succinogenes produces not only succinic acid but also acetic and formic acids as byproducts, additional processing is required to purify the succinic acid from the other organic acids, as well as from salts and other contaminant components present in the starting hydrolysate. The design basis for the succinic acid purification operation is based on recent patents by 85 Myriant and Novasep for the recovery of succinic acid. The whole broth is first filtered through an ultrafiltration step to remove cells and large insoluble particles from the broth. These recovered solids are sent to the boiler. The cleaned broth is then introduced to a simulated moving bed unit (SMB-ICX), which converts the sodium succinate to succinic acid via ion exchange. The succinic acid rich stream is then further purified through nano-filtration to help remove any residual solids and impurities that might color the succinic acid crystals. The dilute succinic acid stream is then further concentrated in an evaporation process and then crystallized. The succinic acid crystals are recovered via a centrifugation unit, dried via a belt conveyor, and sent to storage. Approximately 90% of the mother liquor from the crystallizer, which contains nearly all of the unconverted sugar, is recycled back to the fermentation reactor for further conversion. The overall succinic acid recovery efficiency is set to 82% with the bulk of the succinic acid losses occurring in the crystallization process. Economic Analysis Details. Utilizing the conceptual process design described above, Aspen Plus process models were developed to estimate key material and energy flows in the process. The overall variable and fixed operating costs, as well as the capital costs, were then estimated for the integrated design and the calculated flow rates. The capital costs, which are outlined in Table S3 and S4, are based on vendor designs and cost estimates for the major processing equipment. The details of these equipment 22, 41 designs have been published in the Humbird et al. and Davis et al. design reports. The cost associated with the recovery and purification stages in the design for chemicals production was taken from recent design estimates from external vendors. Raw material prices were drawn from a broad range of sources as reviewed in Table S5. A discounted cash flow analysis was used to determine the overall minimum fuel selling price (MFSP) of the process under a specific set of financial assumptions. The financial assumptions used to calculate the MFSP are reviewed in Table 3, based on a th 41, 86-89 mature n -plant and consistent with prior published work. All coproducts were treated using a market-value allocation approach in a consistent manner with previously published studies, in which a credit is accounted for in the operating costs and 41, 86-89 is allocated based on the coproduct market value. On a relative mass yield basis, a higher yield of the chemical coproduct is observed relative to the fuel yield for a number of the cases considered in this study. Alternative strategies could be utilized to estimate the cost of the chemical coproduct rather than the fuel, if the fuel is assigned a specific price. However, the primary focus of this study is to understand the cost of fuel production and this evaluation is outside of the scope of this study.


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Table 3: Financial Assumptions and Design Basis Plant life Plant throughput Cost year dollar Capacity Factor Discount rate General plant depreciation General plant recovery period Steam plant depreciation Steam plant recovery period Federal tax rate Financing Loan terms Construction period First 12 months’ expenditures Next 12 months’ expenditures Last 12 months’ expenditures Working capital Start-up time Revenues during start-up Variable costs during start-up Fixed costs during start-up

30 years 2000 dry metric tons/day biomass 2011$s 90% 10% 200% declining balance (DB) 7 years 150% DB 20 years 35% 40% equity 10-year loan at 8% APR 3 years 8% 60% 32% 5% of fixed capital investment 6 months 50% 75% 100%

RESULTS Fuels-only base case economics Utilizing the design and economic basis described above, an MFSP for the projected “improved” cases was determined for both process pathways and is summarized in Table 4. The overall MFSP for the fuel-only base case is $9.55/GGE with a total capital investment of $920 MM. The cost breakdown for the major process areas is provided in Figure 2A. Roughly 39% of the capital cost is attributed to the biological conversion section, which requires 97 bubble column reactors to process the full hydrolysate stream based on the projected lipid productivity of 0.4 g/L/hr. This translates to intensive capital expenditures for this section of the process and leads to the greatest contribution to the fuel price, accounting for 31% of the MFSP. The annual total variable cost of the process is $84 MM/year. Biomass feedstock contributes the most to this value at $57.5 MM/year, accounting for roughly 25% of the MFSP. The second largest raw material cost occurs in the lipid extraction section, in which the demand for sulfuric acid, hexane, and ammonia translates to an annual cost of $8.2 MM/year or 3.5% of the MFSP, while the hydrogen required for fuel finishing accounts for 1% of the MFSP. Enzymes represent a significant driver accounting for 5% of the MFSP, even at the low loadings of 10 mg/g of cellulose projected in the “improved” design. The total variable cost also includes a credit of $6 MM/yr associated with the sale of excess electricity to the grid. The primary power consumer in the process is the bioconversion section, which requires 36% of total biorefinery electricity demand, attributed to aeration requirements. Overall fixed costs, which are partially a function of capital costs, are $21.2 MM/yr. Economics of coproduction of fuels and chemicals When the base case is modified and the C5-rich stream is separated from the unconverted biomass prior to enzymatic hydrolysis and used to produce chemicals rather than fuel, the overall MFSP of the process is decreased by 45% to $5.28/GGE in the projected “improved” case. Since this modified design diverts some of the sugar stream originally available for fuel production, the overall fuel production rates are reduced from 24 MM GGE/yr in the baseline case to 15 MM GGE/yr in the modified design. The addition of the C5 upgrading strategy also results in an increase in the overall capital cost of the process, with the installed cost of this added process train being just below $225 MM or 32% of the overall capital cost of the facility. There are two major areas that dominate this additional capital cost. Recovery and purification of the succinic acid product accounts for $134 MM, while the additional solid-liquid separation processes add $45 MM and account for 21% of the capital costs of the C5 upgrading section. The remaining 15% of the cost is the result of the continuous biological reactor design and the auxiliary equipment required for succinic acid fermentation. These additional costs of the C5 upgrading section are partially offset by reduced overall throughput in the C6 fuel train, which decreases the number of bubble column reactors to 60 and reduces the installed capital cost of the biological upgrading section by $30 MM. In total, however, there remains a 30% increase in the total facility capital investment to $1,216 MM for this process scenario. The addition of the C5 upgrading train also increases the variable operating cost of the process by 50% when excluding any coproduct credits. The process requires $42 MM/yr for raw materials required for product separations and purification with nearly 90% of this total cost being attributed to regeneration of the resin (sulfuric acid and caustic) utilized in the recovery design and the caustic required in biological conversion. Beyond this, the process now requires an electricity import of 11 MW at a cost of $6.1 MM/yr due to the introduction of a recovery and purification process that is power intensive demanding 990 kWh/ton of


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succinic acid produced. This import is on the same order of magnitude as the export estimated in the baseline design. Due to the higher capital cost and increased complexity of the modified design, the overall fixed operating cost of this scenario was $27 MM/yr. Table 4: Comparison of economics for two scenarios considered (“improved” projection cases) Base Case Model Coproduction of Fuels (Fuels only) and Chemicals MFSP ($/GGE) $9.55 $5.28 Annual Fuel Yield (MM GGE/yr) 24.1 15.0 Chemical Coproduct Yield (MM tons/yr) 0 0.127 Total Capital Investment (MM$s) 921 1,216 Total Variable Operating Cost (MM$/yr) 89 135 Coproduct credit (MM$/yr)* 6 247 Total Fixed Operating Cost (MM$/yr) 21 27 * The coproduct in the fuels only model is electricity, while for the fuels/chemicals case it is succinic acid.

(B) Coproduction of fuels and chemicals

(A) Production of fuels only Feedstock and handling Pretreatment and conditioning Hydrolysis and bioconversion Cellulase enzyme Product recovery and upgrading Wastewater treatment Storage Boiler / turbogenerator Utilities C5 train -2

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Figure 2: Cost breakout by process area for two scenarios considered (“improved” projection cases). (A) Base case fuel-only design; (B) Coproduction of fuels and chemicals design. Sensitivity analysis for coproduction of fuels and chemicals A number of the underlying process design parameters projected in the “improved” case have a significant impact on the economics of the overall integrated design. As highlighted in the single point sensitivity analyses summarized in Figure 3 for the coproduction of fuels and chemicals, cost drivers exist throughout the process design. Therefore it is critical to understand the impact each of these drivers have on the integrated process to improve the MFSP and to prioritize further research. In particular, sugar production and conversion are major drivers throughout the process. In the biomass pretreatment section, if xylose yields could be increased from the target value of 78.5% to 80%, the MFSP would be reduced by $0.15/GGE due to the higher amount of xylose available for the production of succinic acid. Since the succinic acid biological conversion can utilize minor sugars, an increase in arabinose production by 5% would further decrease the MFSP by $0.06. In the enzymatic hydrolysis step, conversion to glucose also has a major impact; an increase in glucose yield from 90 to 95% reduces the MFSP by $0.23/GGE while a decrease of glucose yield to 86% results in an increase to the MFSP by $0.11/GGE. When considering conversion efficiency of the sugars to the desired product, glucose conversion is a major driver in the lipid production train as lowering the yield by 20% would increase the MFSP by $0.49/GGE. In comparison, a similar reduction in xylose conversion in the lipid train only increases the MFSP by $0.05/GGE due to significantly lower levels of xylose available for conversion in this train. In the succinic acid train, decreasing the conversion of xylose by 20% would increase the MFSP by $0.75/GGE. This change is much higher than the impact from glucose on lipid production since succinic acid is substantially more


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valuable than the fuel on a mass basis. Glucose conversion in the succinic acid train also impacts the MFSP, despite the lower level of glucose in this train, with a 20% reduction increasing the MFSP by $0.23/GGE. Although the minor sugar concentration is low, a reduction in conversion still has a significant economic impact. For example, a reduction in arabinose conversion by 20% increases the MFSP by $0.11/GGE. Efficient recovery of the final products, as well as the sugar intermediates, at the lowest cost possible is also a key driver in the overall cost of the process. For the case of lipid recovery, the cost of extraction and recovery accounts for roughly 20% of the overall MFSP. If the lipids could hypothetically be recovered from cells triggered to auto-lyse, the extraction and recovery process could be eliminated and, assuming consistent recovery efficiency with the standard extraction basis, the MFSP would be lowered significantly by $1.10/GGE. Alternatively, making use of the extraction operation for lipid recovery, even small deviations from the target lipid extraction efficiency translate to significant MFSP impacts, e.g. $0.43/GGE attributed to a 7% deviation from the targeted 90% recovery; this is due to the significant capital and operating expenditures that have been invested throughout the process up to this final step, where it is critical to recover as much product as possible. In the case of recovery and purification of the chemical coproduct, the cost of the recovery section accounts for over 65% of the capital cost of the chemical production train and has a significant impact on the overall MFSP of the process. Indeed, when the capital cost of the coproduct recovery and purification section is modified by 25%, the MFSP changes by $0.68/GGE. These results highlight that both fuel and chemical coproduct recovery processes are paramount to controlling costs for this biorefinery scenario. For recovery of sugar intermediates, the solid-liquid separation following pretreatment is also a key cost driver as it impacts succinic acid yields, while separations after enzymatic hydrolysis is a driver for fuels production. For the separation performed after pretreatment, reducing the efficiency from the “improved” target value of 99% to the experimental baseline of 96% increases the MFSP by $0.36/GGE, whereas 100% separation efficiency reduces the MFSP by $0.11/GGE. As noted previously, this solid/liquid separation step represents ~21% of the C5 upgrading capital cost and if the cost for this operation increased or decreased by 50% (also a measure of filtration capacity) then the MFSP would change by $0.45/GGE. The biggest driver in this section is the wash ratio utilized. If the wash ratio were increased by 50%, then MFSP would be increased by $0.52/GGE due to the higher throughput of water sent to the C5 train and the additional capital required to process and purify the diluted stream. Similar to the solid-liquid separation immediately following pretreatment, for the solid-liquid separation process after enzymatic hydrolysis, the filter costs, wash ratio, and recovery efficiencies are major drivers; however, their impact on costs is varied. The filter cost has the largest impact, as modifying this parameter by 50% will result in a $0.19/GGE impact on MFSP. The impact on MFSP is diminished in this section relative to this parameter for the unit following pretreatment, as the absolute filter capital costs are roughly 60% lower due to the reduced throughput in this portion of the process. Variations in the wash ratio also have less of an impact on costs (variations of 50% result in $0.16/GGE impacts to the MFSP), owing to the use of a concentration unit immediately downstream of this filter step, thus a more dilute hydrolysate stream in this case merely impacts the cost of concentration rather than the entire processing train. Yield impacts are also smaller, such that a 5% reduction in sugar yield increases the MFSP by $0.08/GGE. Finally this section also includes the need for flocculant, which when varying the flocculant loading requirements by 50% affects the MFSP by $0.09/GGE. Raw material prices are also key drivers throughout the process. Changes of 50% in the amount of acid required in pretreatment impacts the MFSP by $0.08/GGE, due to both the change in acid required as well as the change in ammonia required for neutralization. In the enzymatic hydrolysis section, enzyme loading is a more important driver. Increasing the required enzyme loading from 10 mg/g cellulose to the current benchmark value of 12 or even higher to 20 mg/g loading increases the MFSP by $0.12/GGE and $0.78/GGE, respectively.


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Figure 3: Sensitivity analysis results for the coproduction of fuels and chemicals process design (“improved” case, $5.28/GGE MFSP basis). The primary biological metrics dictating economics in the C6 fuel train are achievable lipid productivity and lipid content of the cell. As shown in the tornado plot in Figure 3, productivity has one of the biggest impacts on the MFSP of the entire process, particularly at lower values. Given the large capital costs of the aerobic bioreactors in the C6 train, increasing the productivity from 0.4 to 0.6 g/L-hr would decrease the MFSP by $0.53/GGE while decreasing the productivity to 0.2 g/L-hr would increase the MFSP dramatically by $1.54/GGE. In light of this strong, non-linear sensitivity, the impact of productivity is explored further in Figure 4A, which expands to consider a broader productivity curve. The analysis highlights the asymptotic cost responses, where productivity dramatically influences MFSP below a value of roughly 0.5 g/L-hr and then begins to approach a limit beyond which costs do not significantly improve (roughly 1 g/L/hr, which represents the point at which bioconversion costs no longer dominate the overall system economics). If lipid productivity could be improved to a hypothetical 2 g/L-hr, which is more commensurate with the chemicals production step, the MFSP would decrease to $4.07/GGE. In the case of lipid content, if the lipid content in the cells could theoretically be improved from 70 to 95 wt% (although such a high value is unlikely practical), the MFSP would decrease by 10%. This suggests that for such large improvements in the process, the productivity reduces the MFSP by a greater amount than the lipid content. However, if the focus were on incremental improvements for near-term success, then a 25% improvement in lipid content would achieve a larger reduction in the MFSP than a 25% improvement in productivity. Specifically, a 25% enhancement in lipid content would reduce the MFSP by $0.38/GGE while the same improvement in productivity would reduce the MFSP by $0.30/GGE. It is also important to highlight that the assumed price of the coproduct carries a dramatic effect on the overall MFSP of the process. If the coproduct value were decreased from $0.97/lb to $0.75/lb, then the 0.4 g/L-hr productivity basis would increase the MFSP by more than 70% to $9.04/GGE. In contrast, if the coproduct was valued at a higher price of $1.20/lb, then the MFSP for the 0.4 g/L-hr basis would decrease to $1.42/GGE. Similar trends are seen in the lipid content curve as well (Fig 4B), highlighting the dramatic effect of coproduct price on the MFSP.


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Figure 4: Impact of (A) productivity and (B) lipid content on the MFSP for a range of chemical coproduct prices. One additional scenario was considered for the coproduction process. Since both the C5 and C6 trains can convert C5 and C6 sugars to the desired products, the design was modified to consider splitting the sugar stream after enzymatic hydrolysis. This approach not only allows for the expensive solid-liquid separation process after pretreatment to be eliminated but also allows for flexibility in the amount of sugar that is converted to fuels versus chemicals, as might be dictated by the relative value of each fuel or product or other dynamic constraints imposed on the biorefinery. Further, we estimate the removal of the solid-liquid separations process reduces the cost of the concentrated sugar intermediate price from $0.20/lb of sugar for the split case to $0.18/lb of sugar. The impact of this design modification on the process economics is illustrated in Figure 5. It is observed that increased production of the chemical coproduct relative to the fuel significantly reduces the cost of the process. When 10-20% of the hydrolysate stream is upgraded to chemicals, the coproduct price has a minor impact on the MFSP (e.g., the difference is $0.55/GGE for the lowest and highest value coproduct when 10% of the hydrolysate is split to chemicals production). As the fraction of the hydrolysate upgraded to chemicals increases beyond 30%, then the relative production split and the price of the coproduct have a greater impact on the MFSP of the process.


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Figure 5: Impact of fraction of whole hydrolysate split off to chemical coproduction on the minimum fuel selling price for a range of chemical coproduct prices. DISCUSSION As demonstrated here, the inclusion of co-product opportunities into an integrated process design shows great potential to lower the MFSP of hydrocarbon biofuels produced from lignocellulosic biomass. The target case that includes production of a chemical co-product enables the potential to achieve an MFSP more than $4/GGE (45%) lower than the equivalent design that solely produces fuel, when maintaining the same performance and yield targets for fuel production. Given this key point, and in light of significant cost sensitivities around the projected “improved case” performance parameters presented above, it is important to understand current experimental performance and to prioritize research and development directions. When considering the current benchmark case for production of fuels and chemicals, the MFSP is higher than either the targeted fuelonly design or the fuels/chemicals design. Namely, the estimated benchmark MFSP is $11.37/GGE for the fuel/chemicals coproduction pathway, more than twice the projected cost for the fuels/chemicals co-production “improved” case and 20% higher than the fuel-only “improved” case. Figure 6 illustrates the variation of the cost breakout by process section for all three cases. Comparing the two “improved” cases, the overall cost breakouts are similar for all sections on a $/GGE basis (in fact somewhat higher for the fuel/chemical pathway given a lower fuel yield which increases $/GGE cost contributions). It is clear that the reduction in overall MFSP cost is due to the chemical co-product, which provides for substantial coproduct revenue. When comparing the benchmark and improved cases for the coproduction of fuels and chemicals pathway, given the higher cost of the benchmark case, a number of process improvements are needed moving forward to reduce the overall cost of the design to the projected values. It is promising, however, that there are a number of additional opportunities to further reduce costs. Consistent with established approaches for process design and engineering, the overall economics of a process can be enhanced by understanding the extent of process integration and identifying key approaches for intensification. For the biological conversion to fuels and chemicals considered here, there are myriad potential improvements. One of the primary conclusions of our analysis is that for near-term, incremental improvements, the lipid content is a bigger cost driver than productivity for the process, as long as productivity can remain over at least 0.3 g/L-hr. This may help guide R&D priorities, as it is well-established that productivity and lipid content are inversely related, i.e., operational strategies may be employed to target increased organism lipid content, but typically at the expense of slower production rates. Another demonstrated advantage in the proposed design is that the organisms being pursued are capable of producing valued added products from a wide range of sugars and can tolerate impurities common in hydrolysates. To date, promising results have been reported in the literature for several oleaginous yeast species on biomass-derived hydrolysates, although substrate toxicity 70, 90-101 There are a similar number of opportunities on the “C6 is still a major issue that limits lipid content and productivity. train” that involve enhancing organism performance; for example, Blazeck et al. describe an engineered Y. lipolytica strain that has been observed to produce up to 71% lipids at a reasonable productivity rate of 0.2 g/L-hr in small-scale batch culture 69 conditions. Higher lipid levels approaching 90% have also been observed with variants of Y. lipolytica [31]. Quite interestingly, Qiao et al. recently reported another engineered variant of Y. lipolytica that is able to achieve 84.7% of theoretical lipid yield at 102 55 g/L titers and a productivity of 0.71 g/L-hr. However, the bulk of these studies have been focused on small scale cultivations with glucose, and such results will likely not to directly translate to hydrolysates. Thus, while 90% lipids may be too optimistic today for a realistic commercial-scale system utilizing biomass hydrolysate, even incremental changes to both the lipid content and the productivities would improve the overall economics.


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Beyond near-term organism targets based on intracellular lipid pathways, there are other promising options with the potential to significantly improve process economics. For example, published literature describes an “autolysing” yeast strain currently utilized in the wine industry for its resulting flavor profile, whereby the cell is triggered to lyse and release its intracellular 103-104 contents without the need for a dedicated extraction step. If such an organism could be utilized and/or engineered in this context for lipid accumulation and self-induced release, the resulting cost would drop substantially to $4.18/GGE as shown previously in Figure 3. Furthermore, other organisms capable of product secretion (whether bacterial or otherwise) possess the potential to reduce costs even further. Cost savings brought about by removal of extraction capital and operating costs, combined with improved carbon efficiencies to fuel, suggest significant room for improved economics moving forward. Additionally, we stress that lipid production in oleaginous microbes are merely one of many variations of biological fuel precursor production, and indeed, many strategies are being pursued that target other intermediates. As demonstrated here, detailed TEA in each case will be of significant importance to understand the technology drivers for attaining cost-effective production of any given fuel precursor.

Figure 6: Cost breakout comparison of benchmark and improved cases for coproduction of fuels/chemicals scenario and improved case for fuel-only scenario. The grey bar and reported total value correspond to the MFSP for each case. For the production of chemicals, A. succinogenes is able to tolerate impurities common in hydrolysates, including furfural and 84, 105 HMF at concentrations relevant to biomass hydrolysate without substantial decreases in yield. It readily converts mixed sugar streams of both C6 and C5 sugars common in hydrolysates. The capability of the organism to utilize minor sugars, including arabinose and galactose, as well as to enhance the yield for production when CO2 is included in the conversion, may lead to improved carbon efficiency in the overall process and allow for an additional low cost feed stream and carbon source to further improve the economics of the process. If the productivity of this organism could reach the 4 g/L-hr values reported for 106 engineered E. coli on clean commodity C6 sugars, the MFSP would decrease by $0.10/GGE. It is important to further note that the overall succinic acid yield from these E. coli strains are lower than those demonstrated by A. succinogenes. A major driver in the economics of the C5 chemicals production train is the cost of separating succinic acid from the fermentation broth. The bulk of this cost is not in purification or polishing of the final product, but rather the requirement for handling the succinate salt produced in the biological conversion step. The need to maintain the pH of the fermentation near neutral 14 ACS Paragon Plus Environment

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conditions dramatically increases the complexity and operations for this separation. If the cost associated with only the caustic added to the biological conversion could be completely removed, the MFSP would decrease by over $1/GGE. In part, this could 107-108 be accomplished by developing separation strategies in which the salt is recovered and recycled. Another more promising approach would be to perform the biological upgrading at low pH such that the addition of salt was not required and the separation strategy could be simplified to completely eliminate the need for the SMB unit; approaches such as this are under 109 development in acid-tolerant yeast such as Issatchenkia orientalis. If this design strategy is possible and the yield of the succinic acid remained constant, the overall MFSP would be substantially reduced to $2.32/GGE. If a low pH fermentation would not exhibit as high of a yield as that being projected in this design, however, we estimate that the succinic acid yield could decrease by about 18% and still achieve cost parity with the “improved” coproduction of fuels and chemicals case. Due to this clear potential cost advantage as well as the potential for a more simplified recovery scheme, a number of commercial companies producing organic acids have highlighted on-going R&D to develop low pH fermentation routes. Lastly, we also note that the details of the fermentation and separations technology modeled here are specific to succinic acid production; however, many other products can be produced biologically or catalytically from C5-enriched hydrolysate. It is also important to consider the rationale for coproduct selection and the implications of scale-up in terms of the market and commercial biorefinery economics (here specifically considering succinic acid). In 2009, worldwide consumption of succinic acid 110 was 30,000 tons per year with a projected growth of 5% per year. The current design would almost quadruple the 2009 production level with an estimated production rate of 127,000 tons of succinic acid per year from a single biorefinery, if succinic acid were intended to be the final and only product. However, we stress that the intent for this design strategy is to utilize succinic acid as a representative molecule to demonstrate the potential for coproduct pathways within the framework of an integrated biorefinery, specifically using our research data that demonstrates biological production of succinic acid at encouraging performance levels. Moreover, when viewed as an intermediate, there is a strong potential for a greater demand of succinic acid given that there are numerous known technologies for the production of a wide slate of high value, large market commodity chemicals from succinic acid as illustrated in Figure 7. Based on these potential routes to chemicals alone, the market 60, 111 size for succinic acid has the potential to be well over 2,000,000 tons/year. The concept of improving the overall economics of a biorefinery through the production of high value chemical coproducts from 48, 53, 110, 112-114 This modular a fraction of lignocellulosic hydrolysate can be applied to a wide array of other intermediates. approach will allow for flexibility when developing biorefineries, to avoid overwhelming a single individual commodity chemical market, while supporting the financial feasibility using realistic process and market scenarios for modeling coproduct pathways. As outlined previously, the underlying price of the coproduct can have a large impact on the MFSP as some products will not garner a high enough value to overcome the additional cost in both capital and operating expenses attributed to additional processing steps.

41, 48, 53, 115-116

Figure 7: Known succinic acid derivatives

Beyond the projections and additional improvements specific to fuels and chemicals production discussed here, additional opportunities exist regarding overall process integration to maximize yields to the desired products. The fuels and chemical production strategies outlined can convert both C5 and C6 sugars with a high tolerance to impurities in hydrolysates. Customizing the biomass deconstruction strategies to maximize the desirable sugar intermediates through improvements to 9, 117 118 pretreatment and cellulase efficiency may further improve economics, while ensuring high efficiencies of the solid-liquid separation strategies. Additionally, lignin, acetate, and other components in biomass represent an opportunity for additional 119-123 coproducts.


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CONCLUSION The detailed analysis of an integrated biorefinery performed here presents (a) economic potential for a whole-hydrolysate to hydrocarbon fuels process based on plausible near-term improvements to key conversion parameters, and (b) similar potential for introducing a chemical co-production pathway into the biorefinery schematic. We show that the inclusion of coproducts can dramatically improve the overall economic viability of an integrated process for fuel production, similar to the approach commonly employed in petroleum refining. However, there are underlying challenges as the additional processing steps increase the complexity of the process and result in higher overall capital and operating costs for this integrated design. The overall coproduct value also can play a major role in the economics of the process. Key drivers in the integrated design are tied to achievable yields within the individual biological conversion strategies, product separations, and improving the efficiency in the production and retention of sugars. ACKNOWLEDGEMENTS We thank the US Department of Energy BioEnergy Technologies Office for funding this work through Contract No. DE-AC3608GO28308 at the National Renewable Energy Laboratory. We thank Rick Elander for a critical reading of the manuscript, and Michael Bradfield, Adam Bratis, Rick Elander, Jacob Kruger, Jim McMillan, and Willie Nicol for helpful discussions. The U.S. Government retains and the publisher, by accepting the article for publication, acknowledges that the U.S. Government retains a nonexclusive, paid up, irrevocable, worldwide license to publish or reproduce the published form of this work, or allow others to do so, for U.S. Government purposes. 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For Table of Contents Use Only The techno-economic basis for coproduct manufacturing to enable hydrocarbon fuel production from lignocellulosic biomass Mary J. Biddy*, Ryan Davis, David Humbird, Ling Tao, Nancy Dowe, Michael T. Guarnieri, Jeffrey G. Linger, Eric M. Karp, Davinia Salvachúa, Derek R. Vardon, Gregg T. Beckham* Synopsis: Techno-economic analysis of a hydrocarbon biofuel production plant suggests that manufacturing co-products enables significantly more beneficial biorefinery economics.


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The Techno-Economic Basis for Coproduct Manufacturing To Enable

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