Biomass to Liquid Transportation Fuels via Biological and

Article pubs.acs.org/IECR

Biomass to Liquid Transportation Fuels via Biological and Thermochemical Conversion: Process Synthesis and Global Optimization Strategies Logan R. Matthews,§,†,‡ Alexander M. Niziolek,§,†,‡ Onur Onel,§,†,‡ Neesha Pinnaduwage,§ and Christodoulos A. Floudas*,†,‡ †

Artie McFerrin Department of Chemical Engineering, Texas A&M University, College Station, Texas 77843-3122, United States Texas A&M Energy Institute, Texas A&M University, College Station, Texas 77843-3372, United States § Department of Chemical and Biological Engineering, Princeton University, Princeton, New Jersey 08544, United States ‡

S Supporting Information *

ABSTRACT: Biological conversion of biomass into gasoline, diesel, and kerosene provides an alternative means to meet liquid transportation fuel demand, going beyond the traditional thermochemical methods involving gasification, Fischer−Tropsch conversion, and methanol synthesis. Process synthesis is an ideal methodology for comparing the developing biological technologies with established thermochemical methods through input−output modeling of biorefinery units and inclusion in a superstructure. The resulting model takes the form of a mixed-integer nonlinear optimization problem with full heat, power, and water integration. In the novel superstructure, the MixAlco process for biological conversion is modeled, in which biomass is fermented into carboxylic acid salts which are further upgraded into liquid transportation fuels. The model is solved to global optimality based on the minimization of the cost of liquid transportation fuels production using a branch-and-bound global optimization algorithm to provide upper bound solutions, valid lower bound solutions, and an optimality gap. Case studies analyze the performance of the MixAlco process relative to the existing Fischer−Tropsch and methanol synthesis routes when switchgrass is used as the biomass feedstock at capacities ranging from 1 000 bbl per day to 200 000 bbl per day of gasoline equivalent liquid transportation fuels by energy content. From an economic perspective, the MixAlco process is not competitive with the existing methods, with break-even oil prices from $102.08/bbl to $169.06/bbl. However, a combined MixAlco− thermochemical process is also analyzed, with leftover biomass residue from the fermenter converted into synthesis gas and upgraded to fuels. This proposed process provides break-even oil prices of $79.40/bbl to $144.00/bbl, competitive with and even sometimes outperforming the thermochemical methods for fuel production from biomass. The processes are also analyzed for their greenhouse gas impact, and parametric analysis is conducted on investment cost and switchgrass price parameters to study their impact on the biorefineries.

1. INTRODUCTION The production of liquid transportation fuels from a variety of feedstocks is a heavily studied sector of chemical engineering, for good reason. Petroleum consumption amounted to over 35% of the total energy consumption during 2014 in the United States.1 In this same time period, 71% of petroleum usage went directly to transportation fuels.1 Despite the recent reduction in crude oil prices, petroleum remains a nonrenewable resource that must eventually be replaced in both the American and worldwide economies. Currently, only 4.7% of liquid transportation fuel production originates with renewable resources.1 Thus, the transportation industry is an opportune sector for replacing petroleum consumption, and it is imperative that alternative methods of producing liquid transportation fuels are investigated. One such feedstock for reducing petroleum dependence is biomass, a renewable resource with sizable domestic availability and potential to replace a significant portion of petroleum.2,3 Yet, only 13% of biomass consumption in 2014 occurred in the transportation industry.1 Beyond this, ethanol dominates the biomass-based fuel market despite being an inferior product in comparison to gasoline, diesel, and kerosene.4,5 It is much more © XXXX American Chemical Society

appealing to convert biomass directly to these liquid transportation fuels rather than ethanol, and multiple studies exist to investigate possible conversion methods.4,6−10 Biomass provides special appeal in the potential reduction of greenhouse gas (GHG) emissions. Because biomass consumes carbon dioxide through photosynthesis, biomass-based liquid fuels production will help reduce the 32 billion metric tons of carbon dioxide emissions from the consumption of energy worldwide.1,11,12 The availability and environmental benefits of biomass thus ensures it is an appealing feedstock in both standalone and hybrid refineries, despite its relatively high delivered cost which can range from $4.0-$9.0/MM Btu.13,14 Particularly, coal and natural gas are widely available feedstocks which can be converted to liquid fuels in tandem with biomass with significant GHG benefits. Special Issue: Sustainable Manufacturing Received: September 7, 2015 Revised: November 6, 2015 Accepted: November 16, 2015

A

DOI: 10.1021/acs.iecr.5b03319 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

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Industrial & Engineering Chemistry Research

Figure 1. Biomass feed to the combined process superstructure for thermochemical and biological conversion of biomass. The biological fermentation and thermochemical gasification processes are shown in this figure. The two routes are connected as unreacted biomass residue is sent from the fermenter to the biomass drier for downstream hydrogen separation from synthesis gas.

Process synthesis is a powerful tool for analyzing the options available for fuel production from biomass, with multiple applications already. Process synthesis and integration was utilized to analyze the simultaneous ethanol and food production from corn using gasification of the corn stover and fermentation of grain.23 Production of mixed-alcohols from biomass was studied using a proposed “two-stage” optimization method that used given feedstocks and products to find optimal intermediates and conversion steps.24 Mixed-integer nonlinear optimization (MINLP) models have been developed for many applications including process synthesis of fuels and chemicals production from biomass in one facility through pyrolysis, gasification and Fischer−Tropsch (FT) upgrading, and pyrolysis.25 A bicriteria nonlinear programming model was developed for fast pyrolysis of a poplar feedstock to maximize the net present value of a biorefinery while minimizing the global warming potential of its operations;26 similarly, a multiobjective MINLP was developed for gasification and upgrading through Fischer−Tropsch synthesis.27 Finally, the impact of supply and demand uncertainty was studied for a biorefinery that upgrades black liquor into synthesis gas and eventually to fuels.28 The foundation for the advances in this paper is an MINLP process synthesis superstructure utilizing thermochemical conversion of biomass through gasification and upgrading with either FT or methanol synthesis processes.6−8 This superstructure contains full heat, power, and water integration and can be utilized for hybrid refineries that use coal, biomass, and/or natural gas with complete life-cycle analysis.13,29−31 It was

Biomass conversion techniques are a clear focal point for research in the liquid transportation fuel arena. Technologies for liquid fuels from thermochemically derived synthesis gas are well-established and versatile, able to utilize synthesis gas from natural gas reforming and coal or biomass gasification.13−16 Lignocellulosic biomass conversion is not limited to thermochemical techniques, however. It has been demonstrated that a sequence of catalytic reactions can carry biomass to liquid fuels through intermediates such as furfural or γ-valerolactone.4,9,17 There is also considerable research being conducted into the fermentation of lignocellulosic biomass as a method for biological conversion of biomass into fuels. For instance, lubricant base oils and jet fuels can be produced from fermenting sugar cane and upgrading the product acetone, butanol, and ethanol into fuels catalytically.18 Alternatively, the carboxylate approach begins with biological fermentation of biomass into carboxylic acid salts, which are further upgraded into hydrocarbons through a series of reactions.10,19−21 Beyond this, a considerable amount of experimental research is being conducted to develop consolidated bioprocessing techniques for fuel production; one such study demonstrates the production of isobutanol from cellulosic biomass using a fungal−bacterial consortia.22 Even with these limited examples, the variety of conversion methodologies available prompts the obvious problem of determining which biomass-to-liquid transportation fuel (BTL) process is preferable for new biorefineries in the United States. B


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superstructure components such as hydrogen and oxygen production, wastewater treatment, and heat and power integration; thus, these are categorized into a shared section to delineate the portions of the superstructure specific to each conversion route. The superstructure is shown in Figures 1−13. A color scheme is used to distinguish between each route, as well as to indicate the units whose inclusion in the final solution is a decision of the model. The light blue units exist in both routes and will be included in all solutions; dark blue units can be utilized in both routes but may not be in all solutions or can operate at different conditions. For example, the gasifier in Figure 1 can operate at 900, 1000, or 1100 °C. The light/dark scheme continues for the thermochemical route, whose units are represented in red, and the biological route, shown by green units. Gray arrows demonstrate unit connections that are fixed, while the blue, red, and green arrows are not required to be used in all solutions. Finally, all input and output units, representing feedstocks and products, are shown with purple units. Splitters are shown with a gold color, while mixers are represented in teal. While most mixers and splitters are not labeled, the most important units are explicitly marked in the figures. From Figure 1, it is seen that the two routes join through the biomass drier, to which unreacted biomass residue from the fermenter is sent. Thus, some thermochemical operations are required even for the biological conversion process. The superstructure is described in detail in the following sections. 2.1. Description of Thermochemical Conversion of Biomass. 2.1.1. Biomass Handling and Gasification. The biomass species utilized in the section is switchgrass, which is conservatively assumed to contain 25% moisture by weight when received. The full proximate and ultimate analysis is seen in Table 1.40 The switchgrass is dried to 20 wt % moisture or lower in a

recently expanded for chemicals production, including aromatics32 and olefins,33 with these applications demonstrated through a hybrid refinery that utilizes both biomass and natural gas. To find optimum solutions, the MINLP is solved to global optimality using a rigorous branch-and-bound global optimization strategy.34 The aim of this study is to apply process synthesis and global optimization to a biorefinery that can potentially utilize both biological and thermochemical conversion of biomass. Thus, the MixAlco process, a carboxylate conversion strategy, is modeled for implementation into the existing thermochemical process synthesis superstructure. To our knowledge, this will be the first superstructure that is able to directly compare the relative benefits of thermochemical processes with the MixAlco process for biological conversion of biomass through input−output modeling of each unit; heat, power, and water integration; and a life-cycle analysis. This comparison is crucial for evaluating not only the overall cost of fuels production, but also the relative investment costs, carbon efficiencies, and greenhouse gas benefits on a standardized basis in a global optimization framework. It is also expected that these two routes could be joined together for synergistic benefits; the gasification of unfermented biomass residue is already involved in the MixAlco process for hydrogen generation. The existing process synthesis superstructure uniquely allows this synergy to be realized through potential methanol synthesis or FT processing of synthesis gas. The MixAlco mixed-culture fermentation has been demonstrated for biomass feedstocks including switchgrass,35 municipal solid waste (MSW),36 corn stover,37 and sugar cane bagasse.38 For the purpose of this study, switchgrass is utilized as the biomass feedstock, primarily because of the GHG benefits of soil sequestration and CO2 absorption during photosynthesis.39 In the following sections, the key components of the existing superstructure will be briefly described, and the novel implementation of the MixAlco model is presented in detail. The process synthesis superstructure will demonstrate the relative benefits of (i) biomass gasification to synthesis gas, with or without recycle light gas; (ii) biomass fermentation to carboxylic acids; (iii) synthesis gas conversion via the Fischer− Tropsch process or methanol synthesis; (iv) carboxylic acid degradation to ketones; (v) ketone upgrading to alcohols; (vi) hydrocarbon production via standard upgrading with fractionation or ZSM-5 catalytic conversion; (vii) hydrocarbon production via methanol to gasoline (MTG) or methanol to olefins (MTO) and Mobil’s olefins-to-gasoline/distillate process (MOGD); and (viii) hydrocarbon production via alcohol dehydration and olefin oligomerization. The products from these processes will be gasoline, diesel, or jet-fuel (kerosene), with no constraints on the ratios of these products. Liquified petroleum gas (LPG) and electricity are possible byproducts, depending on the final plant topology. The final MINLP framework will allow for a quantifiable comparison of biological and thermochemical biomass conversion routes, with 24 different case studies at six different plant capacities.

Table 1. Switchgrass Proximate and Ultimate Analysis40 proximate analysis (db, weight %) FCb

HHVc

LHVd

feed switchgrass

25 4.6 79.2 16.2 ultimate analysis (db, weight %)

18636

17360

C

H

N

Cl

S

O

feed switchgrass

46.9

5.85

0.58

0.501

0.11

41.5

moist. (ar)

a

ash

VMa

heating values (kJ kg−1)

VM: volatile matters. bFC: fixed carbon. cHHV: higher heating value. LHV: lower heating value.

d

drier before being fed to a gasifier for conversion to synthesis gas. After the moisture is reduced to 20 wt % or lower, the biomass is lockhopped into the gasifier using compressed CO2. Gasification can occur with or without light gas recycle at temperatures of 900, 1000, or 1100 °C. Two cyclones are used to remove particles from the effluent before it proceeds through a tar cracker to convert any remaining tar. The created synthesis gas is then cleaned before being upgraded into hydrocarbons. These units are shown in blue in Figure 1 because of their usage in the biological route as well, as discussed in section 2.2.2. 2.1.2. Synthesis Gas Treatment. Figure 2 represents the synthesis gas cleaning portion of the superstructure. The synthesis gas composition from gasification can be altered in one of four sour water gas shift (WGS) reactors operating at temperatures between 300 and 600 °C. At 300 °C, the forward water gas shift (fWGS) reaction occurs as in reaction 1.

2. COMBINED BIOLOGICAL AND THERMOCHEMICAL PROCESS SUPERSTRUCTURE DESIGN AND MATHEMATICAL MODEL The thermochemical and biological conversion routes are discussed in this section, with emphasis given to the MixAlco process because the thermochemical superstructure has been described previously.6−8,13 Both of these processes may require C


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Figure 2. Synthesis gascleaning process demonstrated with gas composition altered via water gas shift reactions and various synthesis gas contaminants removed before the synthesis gas is fed downstream to various conversion processes.

H 2O + CO ↔ H 2 + CO2

presence of iron and the required Ribblett ratio of 1.0 for these reactors facilitates the rWGS reaction.41,42 The final two units, developed by Mobil Research and Development,43,44 are slurry phase FT reactors that facilitate the fWGS reaction. These units can produce a nominal or minimal amount of wax. For all FT units, the wax is sent to a wax hydrocracker, while the C1−C30 vapor phase product is sent to upgrading. 2.1.3.2. Fischer−Tropsch Hydrocarbon Upgrading. There are two primary upgrading options for FT hydrocarbons. The first alternative, seen in Figure 5, incorporates a ZSM-5 catalytic reactor which operates at 408 °C and 16 bar.43,44 The effluent of this reactor contains mostly gasoline, which can be separated from the distillate and water through fractionation. The water is sent directly to the wastewater treatment section, while the distillate is hydrotreated to create hydrocarbons for liquid fuels. The gasoline range hydrocarbons are sent to the LPG−gasoline separation section (Figure 8) for product purification. The ZSM5 treatment is not the only option for the FT hydrocarbons, however. A Bechtel design, also referred to as standard upgrading, can be used to make each of the product fuels. Before this, though, the water and oxygenates in the FT effluent are removed and directed to the wastewater treatment plant, as shown in Figure 5. The hydrocarbon recovery system at the heart of standard FT upgrading in Figure 6 receives water-lean hydrocarbons and splits them into C3−C5 hydrocarbons, naphtha, kerosene, distillate, wax, offgas, and wastewater.29,45 These components are further upgraded in the Bechtel design using naphtha reforming, a C4 isomerizer, a C5/C6 isomerizer, a C3/C4/C5 alkylation unit, and a saturated gas plant.45,46 These

(1)

At the higher temperatures, 400, 500, or 600 °C, the reverse water gas shift reaction (rWGS) occurs. Alternatively, the synthesis gas can bypass these WGS reactors. The treated synthesis gas is mixed with the bypassed synthesis gas and proceeds through a dual-capture methanol absorption system, also known as a Rectisol unit. If the sulfur content is insignificant, only CO2 is removed in the Rectisol system. The Rectisol system can also remove acid gas with sulfur content, which must be sent to the Claus sulfur removal system in Figure 3. The captured CO2 can be compressed to 31 bar for use in the refinery, compressed to 150 bar for carbon sequestration, or vented to the atmosphere. The cleaned synthesis gas can then proceed to the hydrocarbon production section to produce liquid fuels via the FT or methanol synthesis routes. 2.1.3. Hydrocarbon Production and Upgrading. 2.1.3.1. Fischer−Tropsch Hydrocarbon Production. A wellestablished technology, the FT process allows for hydrocarbons to be produced with a variety of possible characteristics. Six different FT technologies are incorporated into the superstructure, shown in Figure 4, each operating at 20 bar yet with varying temperatures, catalysts, and end products.13 Cobaltbased catalysts are in two of the FT reactors, in which the WGS reaction is not facilitated. These reactors operate at temperatures of 240 and 320 °C. The per-pass conversion in these cobalt-based reactors is set to 60%. Each of the other four FT reactors involve iron-based catalysts. Two of these reactors operate at temperatures identical to those of the cobalt-based reactors (240 and 320 °C); however, the D


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Figure 3. Process flowsheet of the Claus sulfur recovery system to capture the sulfur naturally found in biomass species.

light gases, the MOGD reactor can produce gasoline and diesel at ratios >100, essentially pure gasoline, or at ratios as low as 0.12. In both reactors, it is assumed that complete conversion of methanol occurs. Both the gasoline range MOGD product and the MTG product are sent to the LPG−gasoline separation section. The distillate from the MOGD product is directed to hydrotreaters to gain the final fuels. 2.1.3.5. LPG−Gasoline Separation. As hydrocarbons are produced in the ZSM-5, MTG, or MOGD reactors, the light gas and gasoline mixtures are sent to a series of columns to separate gasoline from components of liquified petroleum gas, which is sold as a byproduct. The hydrocarbons are sent to knockout columns in Figure 8 which separate CO2 containing light gases from crude hydrocarbons at vapor−liquid equilibrium. These crude hydrocarbons are sent to a deethanizer column, which removes essentially all C1 and C2 hydrocarbons in tandem with a lean oil absorption column.47 While the light hydrocarbons are recycled in the refinery, the heavier hydrocarbon effluent is sent to a stabilizer column to separate C3 and C4 hydrocarbons from the C5+ components which can be blended into gasoline. Lean oil is removed from the C5+ hydrocarbons for use in the absorber column before this stream is sent to the gasoline blender. The C3 and C4 hydrocarbons are fed to an alkylizer unit where it is assumed that they are completely converted to isooctane for blending.47 LPG and isooctane leave this unit and are separated, with the LPG sold as a byproduct and the alkylate blended with the C5+ hydrocarbons as gasoline. 2.2. Biological Conversion through the Carboxylate Platform. The following sections outline the major units in the MixAlco process for converting biomass into hydrocarbon fuels. The approximations used for each unit’s utilities and cost functions are based on data from Granda et al.52 and Pham et

units create hydrocarbons in the ranges necessary for gasoline, diesel, and kerosene. 2.1.3.3. Methanol Synthesis. Rather than proceeding through the FT process, synthesis gas consisting of primarily CO and H2 can be converted to methanol before upgrading into hydrocarbon fuels. This process is illustrated in Figure 7. To create the methanol, synthesis gas is fed to a methanol synthesis reactor operating at 50 bar and 300 °C. In the reactor, the methanol synthesis and WGS reactions are assumed to be in equilibrium. The effluent is flashed at 35 °C, with liquid methanol proceeding downstream and 95% of the unconverted synthesis gas recycled back to the reactor. A purge stream containing the remaining 5% is sent to the light gas recycle loop. This purged gas can be used in the refinery for heat or electricity, as discussed in section 2.3.1. The methanol is again heated to 300 °C to form a gas which is expanded to 5 bar, creating electricity. Finally, after the methanol is recooled to 60 °C, a degasser is used to remove entrained gas while recovering 99.9 wt % of the incoming methanol. While the entrained gases are used as fuel in the refinery, the methanol is sent downstream to be converted to liquid transportation fuels. 2.1.3.4. Methanol Conversion. Methanol created in the methanol synthesis reaction can be converted via two primary routes shown in Figure 7. First, the methanol can be sent to a methanol-to-gasoline (MTG) reactor in which methanol is converted to gasoline range hydrocarbons over a ZSM-5 catalyst at 400 °C and 12.8 bar.47,48 Alternatively, the methanol can be converted to olefins (MTO) via a SAPO-34 catalyst at 375 °C and 1 bar.48,49 These olefins are then further transformed into liquid fuels through the Mobil olefins-to-gasoline/distillate (MOGD) reactor, operating at 400 °C and 1 bar with a ZSM5 catalyst.50,51 These two routes produce different ratios of products; while the MTG reactor can produce only gasoline and E


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Figure 4. Process flowsheet for the six different Fischer−Tropsch reactors used to convert synthesis gas into valuable hydrocarbons and wax which are upgraded into liquid fuels downstream. The reactors vary in catalysts (iron- or cobalt-based), temperatures (240, 260, or 320 °C), amounts of wax produced, and whether a forward (fWGS) or reverse (rWGS) water gas shift reaction is present.

Figure 5. First Fischer−Tropsch product upgrading process flowsheet, showing conversion using a ZSM-5 conversion process and preparation for the standard upgrading process by the removal of oxygenate-rich wastewater.

al.20,53 The downstream process continues from Figure 1 into Figure 9 where carboxylic acid salts are upgraded into liquid fuels.

Table 2 contains all relevant nomenclature for the equations in the following sections used to model these processes. A complete F


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Figure 6. Second Fischer−Tropsch product upgrading flowsheet, representing standard upgrading via the Bechtel design. Water-lean FT hydrocarbons are fed to a hydrocarbon recovery unit, separated, and upgraded into each product. A saturated gas plant handles light gases which are converted into liquified petroleum gas.

Figure 7. Process flowsheet for the methanol synthesis and upgrading process. Synthesis gas is converted into methanol which can be fed to the methanol-to-gasoline or methanol-to-olefins and olefins-to-gasoline/distillate processes. Raw gasoline products from these units are sent to the LPG/ gasoline separation section while diesel is directly produced from the olefins-to-gasoline/distillate process. G


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Figure 8. Process flowsheet for the separation of liquefied petroleum gas from gasoline produced in the Fischer−Tropsch ZSM-5 upgrading or methanol-based processes.

Figure 9. Process flowsheet for the downstream MixAlco process to convert carboxylic acids to hydrocarbon fuels.

calcium carbonate. The fermentation occurs anaerobically in a mixed-culture at 55 °C for 4 weeks;35 methanogen inhibitors are utilized to prevent production of methane.35,52 The mixedculture allows all nonlignin biomass components to be digested, including cellulose, hemicellulose, free sugars, and starch.52 Iodoform is also added to these fermentation systems for methane inhibition.52 It is assumed that all of the cellulosic portions of the biomass are converted during the fermentation, leading to an aqueous fermentation broth after the process is complete. However, there are many undigested residues from the biomass, primarily lignin. The conversion of biomass is modeled as an atom balance derived from cellulosic carbon conversion to

listing of all of the indices, sets, parameters, variables, and constraints used in the superstructure can be found in the Supporting Information. 2.2.1. Biomass Pretreatment and Fermentation. Pretreatment of switchgrass preceeds fermentation to allow for digestion of the cellulosic portions of the lignocelluosic biomass. The length and type of pretreatment differs with the type of biomass used, specifically varying with lignin content; it is recommended that switchgrass is pretreated for 2 h at a temperature between 100 and 120 °C.54 Upon completion of the pretreatment, the biomass is fed to a fermenter to produce carboxlyic acid salts in the presence of H


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∑

Table 2. Key Nomenclature for the Portion of the Mathematical Model Listed in Equations 2−20 label

s ∈ SBIO

(2)

definition

indices a s u specific unit indices FERM MASP BDR MAKH OUTW

∑ MWH O·NuS,FERM,H O ≥ raBW ∑

atoms species process units

2

parameters MWs ARs,a raBL raBW raRL raWL raA cfu,u′,s

2

u

S MWs·NSP _BIO,FERM, s

s ∈ SBio

(3)

Here, the molecular weights (MWSs ) of species s are used along with the molar flow rates (NSu,u′,s) between units u and u′; in the

pretreatment and fermentation unit MixAlco acid separation and purification unit biomass drier MixAlco ketone hydrogenation unit output waste from the MixAlco process

case of these equations, the relevant units include the input biomass splitter (SP_BIO), pretreatment and fermentation (FERM), and input lime for pretreatment (IN_LIME). The ratios of input biomass to lime (raBL) and biomass to water (raBW) are based on Hall et al.35 The following two equations ensure that the lime leaves the system and that the molar ratio of calcium to carboxylates is correct in the carboxylic acid salts leaving the fermenter for the descumming and dewatering unit (MASP).

sets (u,s) ∈ SU (u,u′,s) ∈ SUF s ∈ SBIO s ∈ SFERM s ∈ SAcid u ∈ UIn u ∈ UInv

S S MW Ss ·NSP _BIO,FERM, s = raBL · MWLime· NIN_Lime,FERM,Lime

set of species s in unit u set of species s allowed to flow between units u and u′ set of biomass species (i.e., switchgrass) set of products leaving fermenter set of carboxylic acids produced feedstock inputs to the biorefinery units contributing to the total investment cost

S ∑ NuS,FERM,Lime = ∑ NFERM, u ,Lime u

(4)

u

S NFERM,MASP,Ca = 0.5

∑

S NFERM,MASP, s

(5)

s ∈ SACID

molecular weight of species s atomic ratio of atom a in species s required ratio of input biomass to lime in the pretreatment required ratio of input biomass to water in the fermenter fraction of recovered lignin in the fermenter residue ratio of water to lignin in stream to biomass drier fraction of ash to gasifier after fermentation fraction of carbon in unit u going to unit u′ as species s

Finally, the fermentation products are set via an atom balance on the carbon in the fermenter using the atomic ratio (ARs,a) of atoms (a) in each species and the carbon fractions (cfu,u′,s) of each species leaving the fermenter for the separation unit. The conversions in the ketonization and ketone hydrogenation units are set using a form very similar to eq 6. S AR s , C ·NFERM,MASP, s = cfFERM,MASP, s·

∑

NuS,FERM, s·

u ,FERM, s ′

variables NSu,u′,s

AR s ′ , C

molar flow rate of species s from unit u to unit u′

∀ s ∈ SFERM

(6)

Unreacted biomass residues can be sent to the gasifier for conversion to synthesis gas. To concentrate the carboxylic acid salts, the product must be descummed and eventually dewatered before the carboxlyic acid salts can be upgraded to hydrocarbons. 2.2.2. Descumming and Dewatering. The weight fraction of carboxylic acids in the fermentation broth can range from 5% to 6% by weight, meaning a large amount of water must be removed to concentrate the salts.35,53 While some residue from the fermentation can be sent to the gasifier, other impurities are lost as scum through the addition of flocculant, which must be done before the water is removed. Equation 7 represents the sending of unfermented lignin to the biomass gasifier. Here, raRL represents the ratio of recovered lignin sent to the biomass drier to the amount of lignin in the feed switchgrass that remained unconverted in fermentation. The lignin that is not recovered in eq 7 is lost as part of the fermentation scum.

carboxylic acids. Because of the long cycle time for the batch fermentation, a round-robin process is implemented to ensure an essentially steady-state flow and composition leaving the fermenter.35,52,53 The carboxlyic acid distribution leaving the fermenter consists of 77.14% acetic acid, 3.25% propionic acid, 18.46% butyric acid, 0.59% valeric acid, 0.52% caproic acid, and 0.03% heptanoic acid by weight.36 It is assumed that calcium carbonate, required in fermentation, is recycled after the ketonization reaction35 but that lime is not recoverable. The cost impact of potential lime recycle is discussed briefly in section 3.2. The utility requirements for pretreatment and fermentation are derived from Table 5 of Granda et al.;52 thus, it is assumed that these two processes combined will require 20.8 kWh/tonne biomass of electricity, cooling equivalent to 30.4 m3/tonne biomass of cooling water, and 3.7 GJ/tonne biomass of fuel for heating.52 The mathematical model for fermentation was derived using known fermentation data from a similar feedstock36 which provided consistency with a material balance of the switchgrassbased process.35 The equations for modeling the fermentation are provided in eq 2−6. Equations 2 and 3 constrain the inputs to the pretreatment−fermenter unit to ensure enough lime is present for pretreatment and enough water is present in the fermenter, respectively.

S S MWLignin·NMASP,BDR,Lignin = raRL ∑ MWLignin·NMASP, u ,Lignin u

(7)

Equations 8 and 9 ensure that the proper ratio of water to lignin (raWL) and fraction of ash (raA) leave for the biomass drier. S S MWH2O·NMASP,BDR,H = raWL·MWLignin·NMASP,BDR,Lignin 2O

(8) I


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Figure 10. Process flowsheet for the external light gas recycle loop in which light gas is utilized for heat via fuel combustion, for electricity through gas turbines, or is recycled back to conversion processes after being reformed in an autothermal reactor. S MWASH·NMASP,BDR,ASH = raA

∑

S MWASH·NMASP, u ,ASH

This reaction is implemented using the same general form of eq 6 for the fermenter. The calcium carbonate produced in the reactor can be recycled for use in future fermentations, greatly reducing the cost of fresh calcium carbonate.20 The produced ketones are immediately upgraded to alcohols at 55 bar and 130 °C according to reaction 11.53

MASP, u ,ASH

(9)

While the existing MixAlco process uses lignin gasification for eventual hydrogen production, it is also possible that this lignin may become a source of liquid fuels through the thermochemical route. After flocculant is added to the fermentation broth and the descumming occurs, the water is vaporized using a vapor compression system which yields crystallized calcium carboxylate salts.35,52,53 The salts are then sent downstream to a ketonization unit. The removed scum exits the process as waste, while the water from the fermentation broth can be recycled back to the fermenter to reduce the cost of freshwater required for the MixAlco process. 2.2.3. Ketone and Alcohol Production. The final production of hydrocarbons from carboxylic acid salts proceeds through two intermediate chemicals, namely, ketones and alcohols. Carboxlyic acid salts can be converted to ketones through thermal degradation at a temperature of 430 °C and under vacuum, at 0.04 bar.52,53 The reaction is shown in reaction 10, with R and R′ representing separate hydrocarbon groups. R−COOCaCOO−R′ → R−CO−R′ + CaCO3

R−CO−R′ + H 2 → R−CHOH−R′

(11)

A Raney nickel catalyst is utilized for the isothermal reaction, which takes place in three continually stirred tank reactors. To maximize ketone conversion, hydrogen is fed at 20% excess.53 The hydrogenation reaction is modeled as a stoichiometric reaction that occurs with 98.4% conversion.53 The produced alcohols are fed downstream into a final unit which produces hydrocarbon fuels. 2.2.4. Hydrocarbon Production and Oligomerization. Three reactions occur to produce olefins, oligomerize them, and hydrogenate the double bonds to form alkanes of varying ranges for fuels. In a single unit, with a H-ZSM5 catalyst, the alcohols are first dehydrated in reaction 12 and then oligomerized according to reaction 13.53

(10)

The reaction is modeled using the known carboxylic acid salts produced from switchgrass fermentation and an atom balance around the reactor derived from carbon conversion. The reaction reaches 99.5% conversion with the remaining 0.5% assumed to be lost from the system as waste.20,53 A flash unit separates the ketones from the small amount of noncondensable products.

R−CHOH−CH 2−R″ → R−CHCH−R″ + H 2O

(12)

CmH 2m + CnH 2n → Cm + nH 2(m + n)

(13)

These reactions occur at 300 °C and 3 bar. Finally, the oligomerized olefins can be upgraded to paraffins through saturation with hydrogen in reaction 14.53 CpH 2p + H 2 → CpH 2p + 2 J

(14) DOI: 10.1021/acs.iecr.5b03319 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

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Figure 11. Hydrogen and oxygen production process flowsheet is shown. Either an electrolyzer or pressure swing absorption unit is required to produced hydrogen for the MixAlco process, and an air separation unit is required for oxygen if a PSA unit is utilized.

Figure 12. Process flowsheet for handling product upgrading water, in which a sour stripper and biological digestor can be utilized for the removal of oxygenates and sulfur from the wastewater.

This requires a Raney nickel catalyst. The final hydrocarbons are separated according to their length with the C8−C10 hydrocarbons sent to the gasoline blender, the C11−C13 hydrocarbons

For inclusion in the superstructure, the three reactions can be combined into one stoichiometric reaction for each of the inlet alcohols. Thus, the unit is modeled as a stoichiometric reactor. It

blended with kerosene, and C14+ hydrocarbons blended with diesel.

is assumed that the overall conversion of this combined reaction is 98.4%.53 K


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Figure 13. Superstructure components for the utility wastewater treatment are shown, with integration to the heat and power networks and process cooling. Steam is produced and recycled to necessary units, and freshwater is input in this section if necessary.

2.3. Shared Superstructure Components. While each route contains unique components, the superstructure also includes units which could be required by either process. These units are categorized into light gas handling, hydrogen and oxygen production, and wastewater treatment. 2.3.1. Light Gas Handling. Light gases within the facility can be handled in one of two ways in Figure 10. The first is primarily relevant to the thermochemical route, as an internal gas loop recycles light gas back to units which can convert them to valuable products; namely, gas may be recycled to the methanol synthesis, FT, or gasifying units. While some gas must eventually be purged to prevent a buildup, the internal recycle loop can increase the overall carbon conversion of the process. The external gas loop handles the remaining light gas. These light gases are utilized in three possible ways. The light gases can be sent to a fuel combustor to provide heat for process units in the refinery. Likewise, a gas turbine could be utilized to create electricity for the plant. In either of these cases, the effluents are cooled to 35 °C and sent to a knockout unit to remove water. The effluent could either be vented or sent to a CO2 recovery unit, if necessary, for GHG emission purposes. Finally, an autothermal reactor may create synthesis gas which is recycled back into the hydrocarbon production sections. 2.3.2. Hydrogen and Oxygen Production. Hydrogen is a crucial requirement in both conversion pathways. For the biological route, it is necessary for the ketone hydrogenation and oligomerization reactions. In the thermochemical pathway, it is required in the methanol synthesis reaction, among many other units. Thus, as seen in Figure 11, hydrogen is provided internally

for the plant in one of two ways. A pressure swing adsorption (PSA) unit can be utilized to recover hydrogen from synthesis gas, or an electrolyzer may provide both hydrogen and oxygen from water. In the thermochemical route, where both hydrogen and oxygen are required, the presence of a PSA for hydrogen would necessitate an air separation unit to provide oxygen. 2.3.3. Wastewater Treatment. Wastewater is treated to recycle water back into the refinery or safely remove the water from the refinery. Steam is also produced in this section for process units. A sour stripper and biological digestor can be utilized to remove oxygenates, H2S, and NH3 from the streams shown in Figure 12. The sulfur-rich streams from the biological digestor can be sent to either the fuel combustor or the Claus combustor. The water utility needs of the plant are supplied through a cooling tower for cooling water and boiler system for steam, as seen in Figure 13. 2.4. Capital Costs. The total overnight capital for the BTL refinery is calculated beginning with the total direct costs (TDC) for each unit ⎛ S ⎞sf TDC = (1 + BOP) ·C0·⎜ ⎟ ⎝ S0 ⎠

(15)

where C0 represents a unit’s base cost, S0 the base capacity of the unit, and S the actual capacity; the direct costs scale with scaling factor sf. The balance of plant (BOP) percentage is estimated at 20% for these biorefineries and incorporates costs attributed to site preparation, utility plants, etc. Using the chemical engineering plant cost index, the prices are converted to 2014 dollars. L


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Industrial & Engineering Chemistry Research Table 3. Key Investment Cost Parameters (2014 $) for Units in the Biorefinery unit

C0 (MM$)

S0

SMAX

units

scale basis

sf

biomass drier biomass gasifier Fischer−Tropsch unit ZSM-5 reactor ZSM-5 upgrading methanol synthesis methanol-to-gasoline mobil olefins-to-gasoline/distillate pretreatment and fermentation descumming and dewatering ketonization ketone hydrogenation dehydration and oligomerization

14.12 54.01 43.74 20.23 12.92 11.35 9.22 36.46 9.13 20.06 18.42 12.83 26.74

17.944 26.250 54.975 23.786 4.234 35.647 10.628 10.628 11.111 111.111 5.60 2.944 2.983

33.333 33.333 60.0 100.0 100.0 210.0 42.5 100 − − − − −

kg/s kg/s kg/s kg/s kg/s kg/s kg/s kg/s kg/s kg/s kg/s kg/s kg/s

biomass feed biomass feed feed feed feed feed feed feed biomass feed feed feed feed feed

0.77 0.70 0.72 0.72 0.67 0.60 0.65 0.65 0.54 0.69 0.56 0.56 0.56

which is the familiar form seen in Baliban et al.13 These levelized investment costs, along with the feedstock and utility costs required, are used to determine the overall cost of the refinery. 2.5. Objective Function. The overall cost for liquid fuels production serves as the objective function for the global optimization algorithm. Each component is levelized by the total lower heating value of the products and summed according to eq 20.

While the descriptive parameters for the majority of the units in this paper have been provided previously,13 key parameters for the case studies provided are listed again in Table 3. The full table of all investment cost parameters can be found in the Supporting Information. The parameters for the MixAlco process are estimated from the technoeconomic analysis conducted by Pham et al.;53 they represent the cost of the specific units plus any auxiliary units required for this unit to function correctly. Indirect costs (IC) for the plant investment are calculated using the sum of the total direct costs to account for expenses such as engineering, startup, royalties, spare parts, and contingencies. This is estimated as 32% of the TDC. Finally, the total overnight capital (TOC) is calculated as the sum of the TDC and IC. The BOP and IC are not directly included for units such as those in the MixAlco process, whose cost parameters are already estimated from the total overnight capital in their respective technoeconomic analyses. To accurately compare costs across each process route and capacity, the overall costs are levelized based on the lower heating value of product produced, in gigajoules. This occurs using the method described by Kreutz et al.55 Capital charges (CC) are calculated for the refinery as the product of the levelized capital charge rate (LCCR), the interest during construction factor (IDCF), and the overnight capital. CC = LCCR· IDCF·TOC

min

U

LCCR· IDCF·TOCu CAP· Prod· LHVProd

+

TOCu ·OM 365·Prod ·LHVProd

CostUu (20)

The input feedstock costs (CostFS ) represents all inputs required for the process including biomass, butane, lime, calcium carbonate, and water. The cost of electricity (CostEl) is simply the difference between the electricity required for the process and the electricity produced by the process; it is possible for the electricity “cost” to be negative such that electricity is produced and sold as a byproduct. This is also true of LPG byproduct sales (CostLPG). Carbon sequestration is also made available if needed, although it is very unlikely that this would be required for a biomass-based system. Finally, the investment cost for each unit is summed to gain the total overall investment. The minimization of the cost of fuels production provides a convenient avenue for comparison to the modern petroleum product market through the calculation of a break-even oil price (BEOP). The BEOP is the price at which petroleum must be sold in order for the biorefinery process to be competitive, and it is calculated using the refinery margins for gasoline, diesel, and kerosene. Given the cost of fuels production, the break-even revenue from the fuels is known; subtracting the refinery margin from this revenue provides the cost of petroleum, which is converted to a per barrel basis and reported as the BEOP. In 2015, crude oil spot prices have varied between $38/bbl to $62/ bbl from January to August, providing a very competitive baseline for comparison of results.56 The combined MixAlco and thermochemical process synthesis model, along with simultaneous heat, power, and water integration, represents a large-scale nonconvex mixed-integer nonlinear optimization (MINLP) model. With every superstructure component available, the model consists of 32 binary variables, 10 787 continuous variables, 10 881 constraints, and 708 nonconvex terms. Specifically, the nonconvex terms consist of 642 bilinear terms, 1 trilinear term, 3 quadrilinear terms, and

(16)

(17)

(18)

Together, the levelized investment and operating and maintenance costs can be combined into one total levelized cost associated with a unit (CostUu ) CostUu = Cost uINV + CostOM u ⎛ LCCR· IDCF TOCu OM ⎞ ⎟· =⎜ + ⎝ ⎠ CAP 365 Prod ·LHVProd

∑ u ∈ Uinv

where CAP represents the yearly operating capacity, assumed to be 330 days/year. The operating and maintenance (OM) costs are approximated as 5% of the TOC, allowing the calculation of the operating and maintenance cost per unit (CostOM u ), CostOM = u

CostSF + CostEl + CostSeq + CostLPG

u ∈ Uin (u , s) ∈ S

The LCCR value from Kreutz et al. is 14.38%/year, and the IDCF takes a value of 1.076.55 From this, the levelized investment cost for each unit is simply Cost uINV =

∑ ∑

(19) M


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each case study the ratios of gasoline, diesel, and kerosene were completely free such that each topology could produce a distinct set of liquid fuels. This allows the most economical version of each conversion method to be evaluated and compared. The cost parameters for the BTL process are shown in Table 4 and are gathered for the year 2014 whenever possible. The costs

62 concave cost functions. This large-scale MINLP model requires a rigorous branch-and-bound algorithm to find a global optimum that contains a mathematical guarantee of closeness to the optimum solution through the use of lower bounds via linear relaxation. The solution pool feature of CPLEX 12.657 is initially used to generate a set of 150 distinct points at the root node by solving a mixed-integer linear relaxation of the mathematical node. Starting from the root node, each node is branched to create two child nodes. At all other nodes in the branch-andbound tree, a solution pool of 10 points is generated to serve as starting points for solving the original model. At each generated point, binary variables are fixed and CONOPT 358 is used to solve the resulting NLP. If the NLP solve finds an objective value lower than the current upper bound, the current objective value is replaced. After each step, any node with a lower bound within a given tolerance ϵ of the upper bound is removed from the tree, preventing unnecessary calculations. This algorithm continues until all nodes have been solved or a set time limit has been completed. The final solution provides the current upper bound and the value of the lowest lower bound remaining to provide a mathematical guarantee that the actual global optimum is within this optimality gap.

Table 4. Key Economic Parameters Used in the Biorefinery Cost Calculations

a

3. COMPUTATIONAL STUDIES The MixAlco process is directly compared to the existing thermochemical technologies through global optimization of the superstructure with all of the described components present and switchgrass as a biomass feedstock. While the global optimization algorithm will provide solutions, along with their optimality gap, for the entire superstructure as a whole, the methods for biomass conversion are best compared through case studies which separate the different conversion methods. Results from four technologies are presented to gain a comprehensive view of switchgrass conversion. First, the thermochemical route is solved using a Fischer−Tropsch-based route (TC-FT) and a methanol synthesis based approach (TC-M). The MixAlco process (MA) is also solved to optimality, staying true to the process as described in literature without any interaction with the existing thermochemical superstructure beyond gasification of leftover biomass residue for hydrogen production. Thus, runs for the thermochemical process alone are compared to an optimized MixAlco process. Finally, a hybrid MixAlco−thermochemical (MT) route is also compared in case studies. In the MT route, all of the input biomass is first routed to the fermenter as it would be for normal MixAlco process. The MT route then allows the synthesis gas from the unfermented, gasified biomass to be used in the FT or methanol synthesis processes, unlike the MA route which allows for only hydrogen production from synthesis gas. Six different plant capacities were chosen for the trials: 1 thousand barrels per day of gasoline equivalent product by energy content (TBD), 5TBD, 10TBD, 50TBD, 100TBD, and 200TBD. With the four routes considered and the six different capacities, 24 total case studies are presented. A minimum GHG reduction of 50% was included in the superstructure based on a life-cycle analysis for the process, but this did not factor into any topology because of the overwhelmingly beneficial GHG reduction due to the use of switchgrass as a feedstock. While the superstructure has the capability of producing a variety of ratios of liquid fuels using the thermochemical process or a hybrid MixAlco−thermochemical process, the current nature of the fermentation route and downstream reactors limits the capabilities in the superstructure of varying product ratios for a pure MixAlco process. Thus, in

item

cost

ref

switchgrass propanes butanes freshwater CO2 TS&Ma electricity calcium carbonate lime flocculant iodoform MA waste disposal

$139.97/dry metric ton $1.40/gal $1.84/gal $0.50/metric ton $5.00/metric ton $0.07/kWh $176.64/metric ton $121.10/metric ton $991.00/ton $25/kg $18/ton

13 59 60 13 13 61 62 20 20 20

TS&M: transportation, storage, and monitoring.

for each feedstock includes delivery to the plant gate, and the byproducts are assumed to be sold from the plant gate. For CO2 sequestration, the costs for CO2 transportation, storage, and monitoring are listed in Table 4. The costs for CO2 capture and compression are inherently included in the investment costs of the BTL refinery if sequestration is used. As with the mandated GHG emissions reduction level, CO2 sequestration proves to be unnecessary for a BTL refinery. Each case study was evaluated on the Ada supercomputing cluster at Texas A&M University using GAMS 24.3.3, IBM CPLEX 12.6.0.1, and CONOPT version 3.16C. The branch-andbound global optimization algorithm was allowed to execute for up to 100 CPU hours or 2500 nodes in the branch-and-bound algorithm. The final process topology provides information necessary for completing the heat and power integration for the refinery.30,31 The utility costs are already calculated through the process synthesis superstructure, but two smaller optimization problems are required for gaining the minimum investment costs for the heat exchanger network. The heat engine operating conditions and working fluid flow rates, the amount of electricity produced by the engines, the cooling water requirements for the engines, and the location of pinch points are all provided from the optimum topology. These are then used to identify the subnetworks which must be solved to find the minimum number of matches in the heat exchanger network.31,63 With the minimum number of matches known, the heat exchanger topology with the lowest annualized investment cost is calculated using the superstructure methodology described by Elia et al.30 Finally, the heat exchanger network investment cost is added to the existing investment costs from the optimum superstructure topology to gain the final investment, operating and maintenance, and overall costs of the biorefinery. The results for the refineries are broken into multiple segments ranging from the overall process topology considerations, investment costs, overall cost contributions, energy balances, carbon balances, and GHG emissions. These are displayed in Tables 5−9. 3.1. Optimal Process Topologies. The major decisions for a general biorefinery include the temperature of the gasifier; the N


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Industrial & Engineering Chemistry Research Table 5. Notable Topology Considerations for Each Biorefinerya 1TBD gasifier temperature (°C) fWGS/rWGS tempeprature (°C) Min Wax FT Nom Wax FT FT upgrading MTG usage MOGD usage sequestration usage GT usage MixAlco Usage

gasifier temperature (°C) fWGS/rWGS temperature (°C) Min Wax FT Nom Wax FT FT upgrading MTG usage MOGD usage sequestration usage GT usage MixAlco Usage

MA

TC-FT

900 − − − − − − − − Y

1000 − Mobil − ZSM-5 − − − − − 50TBD

5TBD

10TBD

TC-M

MT

MA

TC-FT

TC-M

900 300 − − − − Y − − −

900 300 − − − − Y − − Y

900 300 − − − − − − − Y

1100 − Mobil − ZSM-5 − − − − − 100TBD

MT

MA

TC-FT

900 300 − − − − Y − − −

900 300 − − − − Y − − Y

900 300 − − − − − − − Y

1100 900 300 300 Mobil − − − ZSM-5 − − − − Y − − − − − − 200TBD

TC-M

MT 900 300 − − − Y − − − Y

MA

TC-FT

TC-M

MT

MA

TC-FT

TC-M

MT

MA

TC-FT

TC-M

MT

900 − − − − − − − − Y

1100 − Mobil − ZSM-5 − − − Y −

900 300 − − − − Y − − −

900 300 − − − − Y − − Y

900 300 − − − − − − − Y

1100 − Mobil − ZSM-5 − − − − −

900 300 − − − − Y − − −

900 300 − − − − Y − − Y

900 300 − − − − − − − Y

1100 − Mobil − ZSM-5 − − − − −

900 300 − − − − Y − − −

900 300 − − − − Y − − Y

The gasifiers could operate at 900, 1000, or 1100 °C. The water gas shift reactors in the synthesis gas cleanning section could operate at 300 °C for the fWGS reaction, or at 400, 500, or 600 °C for the rWGS reaction. Note, these reactors differ from the WGS unit prior to the PSA unit. The FT reactors produce minimum (Min Wax) or nominal (Nom Wax) amounts of wax and can operate with either tradtional iron or cobalt catalysts, or as a medium temperature Mobil unit. FT upgrading could occur using a ZSM-5 reactor or through standard Bechtel upgrading. The presence of MTG, MOGD, sequestration, gas turbine, or MixAlco usage is also noted. a

higher-temperature gasifier presumably for the electricity benefits of the excess waste heat. For the TC-FT, TC-M, and MT trials, the hydrocarbon production and upgrading processes are variable; the MT trials must decide whether to use FT and/or methanol synthesis routes to convert synthesis gas to fuels. Yet even in these two thermochemical subcategories, there exist six possible FT units and two methanol conversion routes. Table 5 provides the relevant reactors for each optimum topology. In all of the FT trials, a Mobil iron-based fWGS minimal wax FT reactor is selected. The CO2 created from the fWGS reaction is not noticeably detrimental to the GHG emissions benefits due to switchgrass usage as a feedstock and lower CO2 production in a 1100 °C gasifier. A ZSM-5 upgrading scheme is included in all TC-FT trials to produce primarily gasoline range hydrocarbons with some diesel; the small amounts of wax from the FT units is converted into fuels using a wax hydrocracker and reforming. Unsurprisingly, due to the free product distribution and the production requirements, gasoline was the only product produced in the TC-M route. Surprisingly, however, an MOGD unit was chosen over an MTG unit as the methanol conversion reactor. This is linked directly to the efficiency of methanol conversion directly to gasoline in the MOGD reactor relative to an MTG process, where LPG is produced as a byproduct. With the high switchgrass feedstock cost of $139.97/ dry metric ton, it was optimal for all methanol to go directly to gasoline rather than to a combination of LPG and gasoline. Essentially, the selling of LPG as a byproduct does not offset the extra biomass cost required to reach the barrel per day production requirement using the MTG route. The methanol synthesis and MOGD process is the primary conversion route for

temperature of the water gas shift reactor, if used; the Fischer− Tropsch units used, if any; the methanol conversion reactors used, if any; the use of carbon sequestration; the use of a gas turbine for electricity generation; and the use of the MixAlco process. Each of these considerations is listed for all of the 24 case studies in Table 5. For a pure MixAlco process, the decisions are relatively limited; the fuel production will follow the typical MixAlco path through biomass to carboxylic acids salts, then ketones, then alcohols, and finally hydrocarbon fuels. The unconverted biomass from the fermenter will proceed to the gasifier, and synthesis gas is used to produce hydrogen. Thus, the only major topological decision for the MixAlco process is the gasifier temperature, a relevant decision for all conversion routes. The gasifier temperature could be set at one of three values (900, 1000, or 1100 °C). Across all MixAlco trials, the gasifier operated at 900 °C; this temperature was also utilized for the TCM and MT case studies. However, in nearly all of the TC-FT trials, a gasifier operating at 1100 °C was used. At the lowest capacity, a 1000 °C gasifier was chosen for the TC-FT trial. A low-temperature gasifier utilizes less oxidant and thus is beneficial if CO2 consumption is not required. The lower temperature has disadvantages regarding CO2 conversion and waste heat production, as this waste heat could be used for electricity generation. However, for a pure MixAlco process, the excess synthesis gas from hydrogen production is sent to the fuel combustor, providing plenty of heat for electricity generation in the heat and power network. There is also no need for the CO2 conversion benefits from higher-temperature gasification due to the GHG benefits of biomass as a feedstock. This latter fact explains why the vast majority of the processes operate with the lowest-temperature gasifier. The FT case studies utilized the O


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Table 6. Overall Contributions to the Cost of Liquid Fuels Production, Normalized by the Heating Value of the Fuels Produced 1TBD MA biomass butane water investment CO2 TS&Ma MA variable costsb O&Mc electricity LPG total ($/GJ) total ($/bbl) lower bound ($/GJ) optimality gap (%)

biomass butane water investment CO2 TS&Ma MA variable costsb OMc electricity LPG total ($/GJ) total ($/bbl) lower bound ($/GJ) optimality gap (%)

15.24 − 0.09 11.77 − 3.72 3.11 −0.87 − 33.05 169.06 32.06 3.01

T-FT

5TBD

TC-M

11.97 12.98 0.05 − 0.01 0.01 14.20 13.60 − − − − 3.75 3.59 0.09 −1.00 −0.06 − 30.01 29.18 152.29 147.66 23.94 25.52 20.23 12.56 50TBD

MT

MA

TC-FT

11.38 − 0.05 11.25 − 2.78 2.97 0.09 − 28.52 144.00 25.94 9.06

15.24 − 0.09 6.57 − 3.72 1.73 −1.00 − 26.34 132.11 24.45 7.17

TC-M

10TBD MT

MA

12.14 12.98 0.05 − 0.01 0.01 8.52 7.84 − − − − 2.25 2.07 −0.16 −1.00 −0.06 − 22.76 21.90 112.36 107.67 21.26 19.90 6.59 9.10 100TBD

11.38 − 0.05 6.08 − 2.78 1.61 0.09 − 21.99 108.15 20.47 6.89

15.24 − 0.09 5.08 − 3.72 1.34 −1.00 − 24.46 121.78 24.13 1.34

TC-FT

TC-M

MT

12.00 12.98 0.05 − 0.01 0.01 7.06 6.50 − − − − 1.86 1.72 −0.04 −1.00 −0.06 − 20.87 20.21 102.01 98.35 17.34 18.33 16.95 9.32 200TBD

11.38 − 0.05 4.71 − 2.78 1.24 0.09 − 20.25 98.60 18.87 6.83

MA

TC-FT

TC-M

MT

MA

TC-FT

TC-M

MT

MA

TC-FT

TC-M

MT

15.24 − 0.09 3.03 − 3.72 0.80 −1.00 − 21.87 107.58 20.55 6.03

11.73 0.12 0.01 5.43 − − 1.43 0.31 −0.09 18.94 91.40 15.55 17.86

12.98 − 0.01 4.76 − − 1.26 −1.00 − 18.01 86.27 16.17 10.17

11.38 − 0.04 2.74 − 2.78 0.72 0.09 − 17.76 84.90 16.38 7.76

15.24 − 0.08 2.54 − 3.72 0.67 −1.00 − 21.24 104.10 20.31 4.36

12.14 0.05 0.01 4.73 − − 1.25 −0.16 −0.06 17.96 86.04 15.11 15.89

12.98 − 0.01 4.38 − − 1.16 −1.00 − 17.53 83.65 16.06 8.35

11.38 − 0.04 2.25 − 2.78 0.59 0.09 − 17.13 81.49 15.99 6.67

15.24 − 0.08 2.25 − 3.72 0.59 −1.00 − 20.87 102.08 20.16 3.42

12.43 0.05 0.01 4.42 − − 1.17 −0.37 −0.06 17.65 84.32 15.39 12.79

12.98 − 0.01 4.05 − − 1.07 −1.00 − 17.10 81.33 15.50 9.36

11.38 − 0.04 1.95 − 2.78 0.51 0.09 − 16.75 79.40 15.85 5.40

TS&M: transportation, storage, and monitoring. bIncludes costs for fermentation and separation such as lime, calcium carbonate, and flocculant. OM: operating and maintenance costs.

a c

turbines were used in only one of the FT trials to produce electricity; in all other cases, light gas was either recycled to units in the process, such as the FT reactors, or sent to a fuel combustor which provided process heat that could also be used for electricity generation via steam production if excess heat is available. 3.2. Comparison of the MixAlco Process to Thermochemical Processes. Across all capacities, the MixAlco process is not competitive with the existing thermochemical processes for converting biomass to fuels. This is immediately evident in the overall cost of liquid transportation fuels production shown in Table 6; the costs range from $17.10/GJ to $29.18/GJ for the thermochemical process across all scales, while the MixAlco process has much higher costs ranging from $20.87/GJ to $33.05/GJ. Even at the largest scale, the MixAlco process requires a break-even oil price (BEOP) of $102.08/bbl. The thermochemical routes provided this BEOP at capacities of 10TBD, and at a maximum capacity the methanol synthesis process provided fuels with a BEOP of $81.33/bbl. This large discrepancy in economic performance can be traced back to the raw material costs for the processes and low carbon efficiency of the MA process. For example, in the 1TBD TC-M case, 42.5% of the carbon entering the thermochemical process leaves as fuel; as a comparison, the 1TBD MA case exhibits an overall carbon conversion of only 35.4%. As seen in Table 8, this majority of this carbon is either vented or lost as part of the fermentation scum. The amount of vented carbon must be

synthesis gas in the MT case studies as well. An MTG process is only used at the 10 TBD capacity; this demonstrates that the MTG and MOGD routes are likely very close in terms of overall costs, and locally optimum solutions for the upper bound of the MT route can be found at comparative levels for both methanol conversion methods. The final observations from the optimum topologies involve the WGS reactor in the synthesis gas cleaning section of the process, as well as the lack of usage of possible sequestration and gas turbines. While there will always be a low-temperature WGS reactor before the PSA unit, an fWGS or rWGS reactor is not necessarily required for synthesis gas in the cleaning section of the flowsheet. Thus, four possible WGS reactors are available in the synthesis gas cleaning section with varying temperatures (300, 400, 500, and 600 °C). At 300 °C, the fWGS reaction occurs, while the higher temperatures facilitate the rWGS reaction. The rWGS reactors are never used in any case study. The fWGS reactor is used in nearly all non-FT case studies. Because a 1100 °C gasifier that produces less CO2 is used in tandem with an FT reactor that facilitates a fWGS reaction, there is no need for the high-temperature fWGS reactor in the synthesis gas cleaning section. However, nearly every other case study utilizes an fWGS reactor in the synthesis gas cleaning section, even the pure MixAlco trials. None of the topologies required sequestration because of the GHG benefits of switchgrass and the cost of sequestration. All topologies were well below the minimum mark of 50% GHG reduction. Gas P


a

Q

H&P: heat and power.

fermentation syngas generation syngas cleanup HC production HC upgrading H2/O2 production H&Pa integration wastewater treatment total (MM$) total ($/bpd)

fermentation syngas generation syngas cleanup HC production HC upgrading H2/O2 production H&Pa integration wastewater treatment total (MM$) total ($/bpd) TC-M − 1051.334 640.916 329.503 406.756 50.041 236.238 92.257 2807.045 56141

TC-FT − 960.241 509.882 379.599 938.001 85.865 244.492 83.168 3201.248 64025

MA

350.398 330.570 194.320 214.315 183.726 136.347 177.223 201.767 1788.666 35773

TC-M − 41.921 39.981 19.064 27.917 5.838 18.327 7.318 160.366 160366

TC-FT − 39.564 34.065 15.280 46.855 8.137 16.271 7.246 167.418 167418 50TBD

MA

22.041 16.435 14.017 23.968 20.524 12.994 12.284 16.526 138.789 138789

1TBD

Table 7. Investment Costs (MM$) for Each Case Study MT

274.517 254.467 178.903 244.143 300.549 96.679 68.575 198.250 1616.083 32322

MT

18.284 13.313 14.278 25.822 28.804 9.786 9.143 13.201 132.631 132631

MA

629.760 616.865 327.096 315.958 270.898 226.443 284.280 322.578 2993.878 29939

MA

62.307 52.695 43.656 59.027 50.583 33.493 38.924 46.294 386.979 77396

TC-FT

− 1848.205 958.073 621.545 1536.259 131.459 343.776 135.965 5575.282 55753

TC-FT

TC-M

TC-M

− 134.567 113.396 51.700 79.483 13.980 48.264 20.720 462.110 92422

− 1961.859 1188.361 614.874 758.482 74.223 425.343 144.765 5167.907 51679

− 127.938 94.403 48.121 140.958 21.960 48.410 20.565 502.355 100471 100TBD

5TBD MT

491.678 474.858 289.765 370.224 456.859 160.237 152.369 254.945 2650.935 26509

MT

51.573 42.663 40.249 64.747 75.321 24.757 22.487 36.799 358.596 71719

MA

1155.584 1151.111 599.975 465.807 399.435 420.061 525.882 584.950 5302.805 26514

MA

97.831 87.113 68.302 87.022 74.581 50.714 61.082 72.071 598.716 59872

TC-FT

TC-M

− 3521.464 1796.762 1159.235 2814.509 304.880 612.631 221.405 10430.886 52154

TC-FT

− 3660.958 2205.419 1147.396 1414.500 118.141 771.913 227.249 9545.576 47728

TC-M

− − 230.251 246.983 147.756 178.086 81.585 81.167 230.960 124.732 32.238 20.452 77.695 82.996 31.412 32.464 831.897 766.880 83190 76688 200TBD

10TBD MT

888.912 886.103 532.070 584.376 747.898 286.545 244.670 422.611 4593.185 22966

MT

80.904 70.510 62.963 96.493 114.116 37.119 35.643 57.255 555.003 55500

Industrial & Engineering Chemistry Research Article


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Industrial & Engineering Chemistry Research Table 8. Overall Carbon Balance (kg/s) for the Biorefineries 1TBD MA biomass butane air calcium carbonate gasoline diesel kerosene LPG fermentation scum unreacted acids unreacted ketones unreacted alcohols vented CO2 seq CO2

biomass butane air calcium carbonate gasoline diesel kerosene LPG fermentation scum unreacted acids unreacted ketones unreacted alcohols vented CO2 seq CO2

TC-FT

3.251 − 0.002 0.021 1.081 0.052 0.027 − 0.569 0.008 0.019 0.019 1.500 −

2.555 0.003 0.002 − 1.086 0.041 0.032 0.004 − − − − 1.398 − 50TBD

5TBD

TC-M

MT

MA

2.769 − 0.003 − 1.179 − − − − − − − 1.593 −

2.430 − 0.001 0.016 1.106 0.039 0.020 − 0.425 0.006 0.014 0.014 0.822 −

16.257 − 0.009 0.105 5.406 0.260 0.133 − 2.847 0.038 0.096 0.094 7.498 −

TC-FT

TC-M

12.958 13.846 0.016 − 0.009 0.013 − − 5.430 5.895 0.204 − 0.159 − 0.021 − − − − − − − − − 7.169 7.964 − − 100TBD

10TBD MT

MA

12.147 − 0.002 0.078 5.530 0.194 0.099 − 2.127 0.029 0.072 0.070 4.107 −

32.515 − 0.018 0.210 10.812 0.519 0.266 − 5.693 0.077 0.192 0.189 14.996 −

TC-FT

TC-M

25.616 27.692 0.031 − 0.017 0.026 − − 10.863 11.790 0.406 − 0.316 − 0.041 − − − − − − − − − 14.037 15.928 − − 200TBD

MT 24.294 − 0.004 0.157 11.059 0.388 0.198 − 4.254 0.057 0.143 0.141 8.214 −

MA

TC-FT

TC-M

MT

MA

TC-FT

TC-M

MT

MA

TC-FT

TC-M

MT

162.574 − 0.091 1.050

125.197 0.378 0.084 −

138.461 − 0.128 −

121.465 − 0.022 0.784

325.149 − 0.183 2.099

259.155 0.310 0.176 −

276.922 − 0.255 −

242.933 − 0.045 1.569

650.298 − 0.366 4.199

530.446 0.616 0.374 −

553.844 − 0.510 −

485.858 − 0.089 3.137

54.059 2.597 1.328 0.000 28.467

49.437 4.674 3.871 0.312 −

58.950 − − − −

55.296 1.940 0.992 − 21.268

108.119 5.193 2.656 − 56.933

108.605 4.072 3.172 0.413 −

117.900 − − − −

110.592 3.880 1.984 − 42.537

216.237 10.387 5.311 − 113.866

217.304 8.092 6.299 0.825 −

235.800 − − − −

221.184 7.760 3.968 − 85.073

0.383 0.958

− −

− −

0.287 0.716

0.767 1.916

− −

− −

0.573 1.432

1.534 3.833

− −

− −

1.146 2.863

0.943

−

−

0.704

1.886

−

−

1.409

3.771

−

−

2.818

74.981 −

67.365 −

79.638 −

41.068 −

149.962 −

143.379 −

159.277 −

82.139 −

299.923 −

298.916 −

318.554 −

164.272 −

reactors required for a Fischer−Tropsch or methanol synthesis process. One potential area of improvement is with lime recycle, which may be possible through a lime kiln.52 In these trials, only calcium carbonate is recycled, but costs could be reduced further if lime is also recycled. To test this impact, trials were conducted with 85% of the lime recycled. With lime recycle, the overall cost is dropped to $30.66/GJ at 1TBD and $18.45/GJ at 200TBD. These prices are more competitive, but are still not as low as the $29.18/GJ and $17.10/GJ from the TC-M process at these respective capacities. 3.3. Comparison of Thermochemical Processes to a MixAlco−Thermochemical Combination. The amount of unconverted biomass leaving the fermenter provides enough synthesis gas to go beyond the hydrogen needs of the MixAlco process, allowing for substantial reduction in the overall costs of liquid fuels production when paired with a typical synthesis gas conversion method for fuel production. This is very clearly seen with a brief cost analysis of the pure and combined fermentationbased processes. With the use of methanol synthesis technologies in the MT route, the costs drop from $33.05/GJ to $28.52/GJ at a 1TBD capacity, and from $20.87/GJ to $16.75/GJ at the 200TBD capacity. The use of synthesis gas from the unreacted biomass reduces the biomass costs from $15.24/GJ in the MA case studies to $11.38/GJ for the MT studies, which is the most

reduced for the MA process to be competitive with the thermochemical routes, which do not lose carbon in a fermentation scum. The overall carbon efficiency would be improved as well if 100% conversion was met in the reactors downstream from the fermenter. Because of the low carbon efficiency, biomass becomes a large cost component for the MixAlco process. At the 1TBD capacity, the biomass costs are over $2/GJ higher in MA than TC. Beyond this, the lime, calcium carbonate, and other variable costs unique to the MixAlco process add to the overall cost. Without improvement to the carbon conversion, these costs will continually push the expenses of the MixAlco process above the thermochemical routes. A potential appeal of the MixAlco process is most evident in its low investment costs, especially at larger scales. These costs are 87% of the TC-MT process at the smallest scale and are only 56% at the largest. This provides great promise to potential investors, given that the technology can be improved to a point where overall carbon conversions increase and the process becomes economically viable. The lack of investment cost is primarily attributed to the biomass conversion technique and the final hydrocarbon production method. Fermentation is much less expensive than gasification, and using parameters from technoanalysis provided by Pham et al.,53 the cost of hydrocarbon production reactors in the MixAlco process scale better than the R


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Industrial & Engineering Chemistry Research Table 9. Overall Emissions Balances (kg CO2 equiv/s) for the Biorefineries 1TBD MA biomass butane lime calcium carbonate gasoline diesel kerosene LPG fermentation scum vented CO2 sequestered CO2 LGHG GHGAF GHGAE GHGI

biomass butane lime calcium carbonate gasoline diesel kerosene LPG fermentation scum vented CO2 sequestered CO2 LGHG GHGAF GHGAE GHGI

−18.79 − 0.50 < 0.01 3.99 0.19 0.10 − 2.55 5.49 − −5.96 5.84 0.29 −0.97

5TBD

TC-FT

TC-M

MT

MA

−14.77 < 0.01 − − 4.01 0.15 0.12 0.02 − 5.12 − −5.35 5.87 −0.03 −0.92 50TBD

−16.01 − − − 4.28 − − − − 5.84 − −5.89 5.84 0.33 −0.96

−14.04 − 0.37 < 0.01 4.06 0.14 0.07 − 1.91 3.01 − −4.47 5.84 −0.03 −0.77

−93.96 − 2.51 0.01 19.95 0.96 0.49 − 12.77 27.47 − −29.80 29.18 1.66 −0.97

TC-FT

TC-M

−74.89 −80.03 < 0.01 − − − − − 20.06 21.39 0.75 − 0.58 − 0.08 − − − 26.27 29.18 − − −27.15 −29.46 29.34 29.18 0.26 1.66 −0.92 −0.96 100TBD

10TBD MT

MA

−70.21 − 1.87 0.01 20.32 0.72 0.37 − 9.54 15.05 − −22.34 29.18 −0.14 −0.77

−187.93 − 5.01 0.03 39.91 1.91 0.98 − 25.54 54.95 − −59.60 58.37 3.32 −0.97

TC-FT

TC-M

−148.05 −160.05 < 0.01 − − − − − 40.13 42.78 1.50 − 1.16 − 0.16 − − − 51.44 58.36 − − −53.67 −58.91 58.68 58.37 0.14 3.32 −0.91 −0.96 200TBD

MT −140.41 − 3.75 0.02 40.64 1.43 0.73 − 19.08 30.10 − −44.67 58.37 −0.29 −0.77

MA

TC-FT

TC-M

MT

MA

TC-FT

TC-M

MT

MA

TC-FT

TC-M

MT

−939.60 − 25.07 0.13 199.54 9.57 4.89 − 127.68

−723.61 0.05 − − 182.52 17.22 14.26 1.18 −

−800.27 − − − 213.90 − − − −

−702.04 − 18.73 0.09 203.18 7.15 3.65 − 95.39

−1879.29 − 50.13 0.25 399.09 19.14 9.78 − 255.36

−1497.84 0.04 − − 401.22 15.00 11.68 1.56 −

−1600.54 − − − 427.81 − − − −

−1404.13 − 37.46 0.19 406.35 14.30 7.31 − 190.79

−3758.58 − 100.26 0.50 798.18 38.28 19.56 − 510.71

−3065.86 0.07 − − 802.79 29.81 23.20 3.11 −

−3201.07 − − − 855.62 − − − −

−2808.11 − 74.91 0.37 812.70 28.60 14.62 − 381.57

274.74 − −297.98 291.83 16.60 −0.97

246.84 − −261.56 294.17 −5.10 −0.90

291.81 − −294.56 291.83 16.60 −0.96

150.48 − −223.36 291.83 −1.44 −0.77

549.49 − −596.05 583.65 33.20 −0.97

525.37 − −542.98 586.76 5.14 −0.92

583.62 − −589.11 583.65 33.20 −0.96

300.97 − −446.77 583.65 −2.89 −0.77

1098.98 − −1192.10 1167.31 66.39 −0.97

1095.29 − −1111.60 1173.50 24.45 −0.93

1167.24 − −1178.21 1167.31 66.39 −0.96

601.92 − −893.41 1167.31 −5.77 −0.77

significant cost reduction. Cost benefits are also seen among the investment costs as well as the MA specific variable costs; with the synthesis gas now being fully utilized, less fermentation is required to meet the production specifications. Most importantly, the combined MT case studies become economically competitive with the thermochemical routes across all capacities when switchgrass is used as a feedstock. The final TC-M and MT costs of liquid fuels production are never separated by more than $0.66/GJ at any capacity. A methanol synthesis and MOGD based process has the lowest upper bound at 5TBD and 10TBD, while a MixAlco process with methanol synthesis and upgrading provides the best solutions at 1TBD and at the highest three capacities. While the pure MA case studies reached only 35.4% conversion of input carbon to liquid fuels, the combined MT trials now reach over 47.6% carbon conversion to fuels, superior to the 42.5% of the TC-M trials and 44.6% in TC-FT trials. The MT trials are slightly more efficient from an energy standpoint as well, as seen in Table S1 of the Supporting Information. Substantial improvements in process efficiency and the vast reduction in investment costs relative to the TC-M route serve as the critical factors in the improved economics. These trials are not substantially preferable in price, however, due to the expensive requirements of the fermenter and separators, including the pretreatment lime, calcium carbonate, and flocculant. Essentially, the reduction in

investment and OM costs relative to the TC-M trials are at least partially offset by the variable costs required for pretreatment, fermentation, and salt concentration from the fermentation broth. According to these case studies, the combined approach for biological and thermochemical conversion of biomass actually reduces the total investment cost required relative to the pure biological or pure thermochemical approaches. The combined MT topology reduces the equipment size needed for the entire MixAlco process, as well as in the hydrogen production section, offsetting the capital required for methanol synthesis processing. Because the majority of the hydrocarbons in the MT process are coming from the fermentation route, the capital requirements for the gasifier and synthesis gas cleaning are vastly reduced relative to the TC processes; in fact, this reduction is nearly 4-fold at the highest capacities. Of all the sections in the biorefinery, the only areas in which an MT topology is more expensive include the fermentation, hydrogen production, and wastewater treatment sections. This naturally follows for fermentation, which does not exist in a TC route. More hydrogen is necessary for the MixAlco process because of ketone hydrogenation and olefin saturation, reactions 11 and 13 respectively, and considerable wastewater is produced in the dehydration reactor of the MixAlco process. As seen in Table 7, the benefits of low investment costs increase for MT case studies as the capacity grows, because of the lower S


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eliminated, but also 77% of the GHG emissions from a comparable petroleum refinery are actually consumed over the lifecycle of a BTL refinery. The pure MA trials all have GHGI of −0.97, which not only is the largest GHG reduction of these trials, but also nearly offsets the entire GHGI emissions of a similar-sized petroleum refinery. For these fermentation-only case studies, the relative inefficiency of the MA process compared to the other trials assists in the GHGI improvements. The extra biomass being consumed as a feedstock more than offsets the extra CO2 that is attributed to MixAlco process waste and the lime and calcium carbonate acquisition. The MT process requires the least biomass and has emissions from lime and calcium carbonate acquisition which the thermochemical process does not require; thus, it has the worst GHGI value of −0.77. As a comparison, the TC-FT process has approximately 15% more GHG reduction on average with a GHGI value of −0.92. The TC-M route has a GHGI of −0.96, nearly equivalent to the MA route. While the GHG reductions of the MT process are still strong relative to a petroleum-based process, the extra GHG improvements of the TC-M process is an added benefit of choosing the more traditional method for biomass conversion to liquid transportation fuels. 3.5. Parametric Analysis of Key Cost Contributors. Two key parameters are considered for the results of the biorefinery case studies. Because the investment cost of the MixAlco equipment is consistently lower than the thermochemical processes, the investment cost is considered. Specifically, the scaling factor of the MixAlco reactors is set to three possible values to ascertain the economical importance of scale-up for this process. The price of switchgrass as a feedstock is also examined; with each process requiring different amounts of biomass to meet the barrel-per-day requirement of the refinery, this parameter has potential for changing the relative feasibility of each process. Because the feedstock price appears in a linear constraint, future work surrounding feedstock pricing uncertainty can utilize robust counterpart optimization.67−73 Each parametric analysis is conducted on the 50TBD scale results, with results at 5TBD shown in the Supporting Information. 3.5.1. Investment Costs. In the existing technoeconomic studies performed on the MixAlco process, the reactors downstream of the fermenters are scaled using a cost exponent of 0.56.20 Accordingly, this value is used as the standard cost exponent for the ketonization, ketone hydrogenation, and hydrocarbon production units of the MixAlco process. These factors are lower, however, than other cost exponents used in hydrocarbon production facilities, specifically for the MTG and MOGD reactors and the FT reactors which use cost exponents of 0.65 and 0.72, respectively.43,44,74 Thus, parametric analysis is conducted on the final topologies found in the aforementioned case studies to measure the impact of investment cost parameters on the overall cost of the process. Figure 14 shows the results of parametric analysis conducted on refineries with a 50TBD production rate. The overall cost of the TC-FT and TC-M routes are shown as a reference; the lack of MixAlco units naturally prevents any change in the overall cost of the themochemical processes. For the pure MixAlco process, in which fermentation is the only hydrocarbon production method, the investment cost increases from $1789 MM at the base exponent to $1933 MM at sf = 0.65 and $2081 MM at sf = 0.72. This is still considerably lower than the overall TC-M investment of $2807 MM, but the rise in investment costs leads to a noticeable increase in the overall cost of fuels production for the MA process. A base overall production cost of $21.87/GJ

scaling factors for downstream MixAlco reactors as listed in Table 3. In light of this, the issue of scale and investment costs is discussed in detail using parametric analysis in section 3.5.1. Ultimately, it is clear that the synergistic benefits of a combined MT process leads to great improvements from a pure MA process and competitive economic performances across all scales relative to thermochemical processes. With BEOPs ranging from $79.40/bbl to $144.00/bbl, the process could even be competitive with petroleum-based processes, depending on the state of the oil market. For this to occur though, the process will have to prove itself to be truly scalable from its current pilot plant performances to a medium or large scale plant evaluated in these case studies. Given the cost similarities between thermochemical and combined topologies, the MixAlco process will also have to demonstrate enough reliability to warrant the added complexity of running multiple conversion technologies, rather than one FT or methanol synthesis based thermochemical plant. The greatly reduced investment costs provide a strong economic incentive for further investigation of a hybrid biological and thermochemical BTL refinery. 3.4. GHG Benefits of the Biorefineries. The GHG benefits of each case study are expected to be significant because the main input is switchgrass. A well-to-wheel lifecycle analysis is performed on the GHG emissions of the BTL refinery. This analysis includes emissions due to the acquisition of switchgrass and other possible feedstocks; the transportation and use of the gasoline, diesel, and kerosene; transportation and sequestration of CO2, if necessary; venting of process emissions; and the capture of CO2 from the atmosphere by switchgrass for soilstorage or photosynthesis. The GREET model is used to calculate emissions associated with the feedstocks and products, with a transportation distance of 50 miles assumed for feedstocks, 100 miles for products, and 50 miles for sequestered CO2.66 Complete combustion of liquid products to CO2 is assumed in the calculations. The amount of CO2 absorbed into switchgrass from the atmosphere for photosynthesis is calculated based on its carbon content; the amount stored in the soil is assumed to be 0.3 g of carbon per gram of as-received biomass over the lifetime of the crop.39 As a conservative estimate, the carbon leaving in the fermentation scum is assumed to completely offset the GHG benefits of an equivalent amount of carbon in the biomass. With these assumptions, the relative benefits of a switchgrassbased refinery can be quantified using a greenhouse gas index (GHGI), which is defined as GHGI =

LGHG GHGAF + GHGAE

(21)

In eq 21, the lifecycle GHG emissions (LGHG) of the BTL refinery are divided by the emissions that would be created from a typical petroleum process. Specifically, the GHG emissions avoided from fuel production (GHGAF) are calculated as being approximately 91.6 kg CO2eq per GJLHV of fuels,64 and GHG emissions avoided from electricity production (GHGAE) are approximated using 101.3 kg CO2eq per GJ electricity.65 If the GHG emissions are equal to those of a typical petroleum process, the GHGI would equal one. A negative GHGI indicates that over the lifecycle of the materials used and produced in the plant, more greenhouse gases are actually consumed than produced. Every case study has negative GHG emissions, as seen in Table 9, because switchgrass is used as a feedstock. In fact, the maximum GHGI value of any case study is −0.77, consistently occurring in the MT trials. This indicates that in the worst case, not only are all net lifecycle emissions from a petroleum refinery T


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studies at a production capacity of 50TBD. It is immediately obvious that the change in feedstock price has a much larger impact on the MixAlco process than on the traditional thermochemical process. Dropping the switchgrass price by $40/dry metric ton leads to a $4.35/GJ reduction in cost. However, this same reduction in switchgrass price produces only a $3.71/GJ reduction in the cost of a TC-M refinery. This follows from the fact that the MixAlco process requires more biomass to produce the same amount of fuels as the thermochemical route. Across all biomass prices considered, the thermochemical processes remain favorable over a pure MixAlco biorefinery, although the gap between the two decreases significantly as the biomass price decreases. The impact of biomass prices is not as drastic for the MT process, but the relevance to the optimum topology is striking. The TC-M process requires more biomass than MT; thus, the gap between the two costs decreases until they are nearly identical ($16.15/GJ) at the switchgrass price of $120/dry metric ton. Furthermore, at $100/dry metric ton and below, TC-M is the least expensive option of all the conversion processes. Thus, a more cost-effective biomass option benefits methanol synthesis the most. The impact of lower biomass prices is most clearly relevant when viewing the adjusted BEOP of each process. At the lowest switchgrass price considered, the BEOPs for MA, TC-M, and MT are $64.77/bbl, $59.05/bbl, $50.75/bbl, and $53.80/bbl, respectively. The latter three BEOPs all fall in the range of crude oil spot prices in the United States during the first six months of 2015. Thus, this variability in switchgrass price is crucial for the economic competitiveness of the process; if the switchgrass price can be lowered to $60/dry metric ton, or fuels can be produced from a different, less expensive biomass feedstock, all of these processes become appealing.

Figure 14. Parametric analysis conducted on the investment cost parameters of MixAlco process reactors in a 50TBD biorefinery with various process topologies.

increases to $22.18/GJ and $22.50/GJ at the higher investment levels in Figure 14, equivalent to an increase from $107.58/bbl to $109.29/bbl and $111.03/bbl. This parametric analysis is even more relevant for the combined MixAlco to thermochemical route, as seen in Figure 14. In fact, with all other parameters constant, the scaling factor used for the MixAlco reactors determines whether MT is favorable over TC-M. When a scaling factor of 0.65 is used, TCM and MT are essentially equivalent in cost; at 0.72, the TC-M route is favorable. The change in investment costs changes the overall cost from $17.76/GJ to $17.99/GJ and $18.23/GJ at scaling factors of 0.65 and 0.72 respectively. Thus, the BEOP is pushed to $86.21/bbl and $87.51/bbl for the latter scaling factors. The change in optimum topology based on these scaling factors gives value to the methanol synthesis route, which is more established. While there is certainly more risk involved in the MT hybrid process, it remains an economically friendly option at all scaling factors; if the process scales as originally suggested, it is the optimum topology considered at the 50TBD scale. 3.5.2. Feedstock Costs. Feedstock cost is another substantial factor in the economic feasibility of a biorefinery, and parametric analysis measures this impact relative to each conversion route and topology. Switchgrass is an expensive feedstock at $139.97/ dry metric ton, and the impact of a price reduction in switchgrass could have major implications for each process. Figure 15 provides insight into this by varying the feedstock price from $60/dry metric ton to $160/dry metric ton for each of the case

4. CONCLUSION Process synthesis and global optimization provide a powerful methodology for evaluating available energy production technologies. Building upon an established thermochemical hybrid feedstock superstructure, the biological conversion process known as MixAlco was modeled and placed into the superstructure to evaluate the economic and environmental competitiveness of the process compared to established thermochemical methods. Switchgrass was utilized as the biomass feedstock as 24 case studies were solved to global optimality using a rigorous branch-and-bound strategy to evaluate a pure MixAlco process compared to two established thermochemical routes and a novel hybrid MixAlco−thermochemical process. While a pure MixAlco process is not yet economically competitive, with BEOPs ranging from $102.08/ bbl to $169.06/bbl, the process is competitive when coupled with a thermochemical method, providing BEOPs between $79.40/ bbl to $144.00/bbl. The combined process even provided lower overall costs of fuel production at high plant capacities when compared to the thermochemical routes. Together, this data gives intrigue to the feasibility of biological conversion routes such as the MixAlco process. Parametric analysis demonstrates the dependence of these processes on feedstock costs, as well as the importance of cost-effective scaling for the downstream MixAlco reactors; without this, thermochemical conversion routes will continue to dominate the BTL economic landscape. Regardless, it is clear that these facilities provide an environmentally friendly methodology for producing liquid fuels from a

Figure 15. Parametric analysis conducted on the switchgrass price in a 50TBD biorefinery with various process topologies. U


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clean, renewable resource which can be competitive in the United States under the right conditions.

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ASSOCIATED CONTENT

S Supporting Information *

The Supporting Information is available free of charge on the ACS Publications website at DOI: 10.1021/acs.iecr.5b03319. More details on the MixAlco process, a parametric analysis at a 5TBD scale, and the complete mathematical model for the process synthesis superstructure (PDF)

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AUTHOR INFORMATION

Corresponding Author

*Phone: 979-458-0253. E-mail: fl[email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS L.R.M. gratefully acknowledges his support from the National Defense Science and Engineering Graduate fellowship. C.A.F. gratefully acknowledges support from the National Science Foundation (CBET-12548540). The authors would also like to acknowledge Professor Mark Holtzapple from Texas A&M University for his assistance involving material balances for the MixAlco process.

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REFERENCES

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