Biomass-Based Production of Benzene, Toluene, and Xylenes via

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Biomass-Based Production of Benzene, Toluene, and Xylenes via Methanol: Process Synthesis and Deterministic Global Optimization Alexander M. Niziolek,§,†,‡ Onur Onel,§,†,‡ Yannis A. Guzman,§,†,‡ and Christodoulos A. Floudas*,†,‡ †

Artie McFerrin Department of Chemical Engineering, Texas A&M University, College Station, Texas 77843, United States Texas A&M Energy Institute, 3372 Texas A&M University, 302D Williams Administration Building, College Station, Texas 77843, United States § Department of Chemical and Biological Engineering, Princeton University, Princeton, New Jersey 08544, United States ‡

S Supporting Information *

ABSTRACT: The pursuit toward an environmentally sustainable energy landscape requires the development of economically competitive renewable processes. Efficient utilization of renewable resources is an important first step toward meeting this goal. To this extent, we introduce a systematic deterministic global optimization-based process synthesis framework that determines the most profitable processes to produce benzene, toluene, and/or xylenes from biomass via methanol. Our framework incorporates several novel, competing, and/or commercial technologies. We quantify the effect that biomass type has on the overall profit of a refinery by investigating forest residues, agricultural residues, and perennial crops as potential feedstocks. A thorough economic analysis, together with material, energy, carbon, and greenhouse gas balances, are provided for every proposed process design. The capability of our proposed approach is illustrated through several case studies that produce varying ratios of p-, o-, and m-xylene across several refinery scales. The most profitable aromatics refineries consistently produce p-xylene, while o-xylene refineries consistently have the lowest required investment costs. The net present values for the biomass to aromatics, BTA, refineries producing 2000 t per day of product are as high as $1200 MM dollars with payback periods less than 10 years.

1. INTRODUCTION The conversion of biomass into more valuable products, such as liquid transportation fuels or electricity, has received significant interest from several United States government agencies, industries, and academia over the past decade. As highlighted in a recent perspective article by Floudas et al., renewable energy research areas, and in particular, biomass based processes, have greatly benefited from multiscale systems engineering components, which include modeling, design, synthesis, simulation, and optimization.1 Two recent reviews have highlighted key contributions in the production of liquid transportation fuels and the supply chain optimization of hybrid energy processes.2,3 A recent perspective article outlined the state of the art approaches and major challenges toward integrated biomass and fossil-fuel systems that can simultaneously provide economical and environmental benefits.4 Biomass has several advantages over the other two feedstocks, natural gas and coal, considered as alternatives to petroleum. Namely, it is a renewable resource that has the ability to absorb atmospheric CO2 during photosynthesis,5−7 thus reducing net life-cycle greenhouse gas emissions. The U.S. Energy Information Administration projects that most of the growth in primary energy consumption until 2040 will occur with an increase in the consumption of natural gas and renewable energy.8 However, in order for biomass to penetrate the large energy market, the processes utilizing it should be profitable and efficient, and it should be harvested sustainably. The former concern has been and is currently being addressed by several studies in the literature that have mainly focused on the production of liquid fuels from biomass via thermochemical, catalytic, or biological routes.9−59 The latter concern has been assessed by the Department of Energy © XXXX American Chemical Society

investigating the technical feasibility of producing one billion tons of biomass in the United States,60,61 and by numerous studies in the literature focusing on the supply chain aspects of biorefineries.62−92 It is remarkable, however, that very few studies have investigated the economic and technical feasibility of producing high-value aromatics from biomass.93−95 Aromatics constitute a large portion (one-third) of the market for commodity petrochemicals,96 and a significant portion (approximately 70%) of the world’s benzene, toluene, and xylenes supply comes from petroleum naphtha.97 Biomass-based aromatics production additionally offsets import requirements for petrochemicals, crude oil, or naphtha, and thus aids in energy independence objectives set by the U.S. government. A stand-alone biomass to aromatics plant has the potential to penetrate the chemical market if it is shown to be profitable. To address this, we propose a process synthesis superstructure composed of several existing or novel process technologies for the thermochemical conversion of biomass into aromatics. The most profitable, optimal process topology is determined using a novel branch-and-bound global optimization framework. The world demand for benzene, p-xylene, o-xylene, and m-xylene was approximately 40, 26, 6, and 0.4 million metric tons per year, respectively, between 2005 and 2008.98 Aromatics serve as midproducts in petroleum refining and have a variety of end uses across several industries.97,99−102 Benzene, for example, is used as a precursor for styrene, phenol, nylon, and aniline production. In 2008, the United States used 5.6 million metric Received: March 15, 2016 Revised: April 30, 2016

A

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Energy & Fuels tons of benzene as a chemical feedstock.101 Toluene is typically blended into unleaded gasoline but is also converted into benzene and xylenes and is used in solvent applications.101 o-Xylene applications include the production of phthalic anhydride, while m-xylene is converted into isophthalic acid. p-Xylene is converted into terephthalic acid and dimethyl terephthalate, which are ultimately used to produce polyethylene terephthalate (PET) fibers, resins, and films.97,103 Despite the fact that aromatics represent a significant portion of the world market for commodity petrochemicals, few studies have considered stand-alone production of aromatics from alternative sources. Lin et al. investigated the production of p-xylene from lignocellulosic biomass and analyzed the economics and life cycle emissions for their proposed processes in a series of papers.93,94,104 Ashraf et al. simulated and developed a process for the methylation of toluene to produce p-xylene using Aspen Plus.105 Several other studies considered the coproduction of liquid transportation fuels and aromatics. Wang et al. used ChemCad to investigate the coproduction of ethanol with aromatics, olefins, and synthetic gasoline and diesel.106 In our previous work, we investigated the coproduction of C6−C8 aromatics and liquid fuels from biomass and natural gas using an optimization-based process synthesis framework and calculated high net present values for these refineries.95 We also recently proposed a stand-alone chemicals facility for the conversion of natural gas into aromatics (GTA).107 Additionally, we have investigated the coproduction of olefins and liquid fuels from biomass and natural gas, as well as stand-alone production of olefins from natural gas.108,109 In order to rigorously compare existing and novel process designs, we include several technologies within a process synthesis superstructure. The methanol-to-aromatics (MTA) reaction has been identified as one promising alternative for the production of aromatics. Several studies investigated this reaction over ZSM-5 and H-ZSM-5 catalysts impregnated with metals, such as Zn-,110−112 La/Zn-,113 Cu-,111,114 Zn/Sn-,115 Ag-,114,116 as well as others.111,114,117 Recently, a pilot plant which utilized this technology and produced 10 ktpa (kiloton per annum) of aromatics was successfully operated by Huadian Coal Industry Group Company in China. Coal was used as the feedstock in this pilot plant because of the lack of oil and natural gas in China, and the company has plans to build a 1 million metric tons per year commercial scale coal to aromatics plant in the upcoming years.118 Additionally, toluene methylation has been identified as an attractive alternative for xylenes production119−125 since the amount of toluene used as a chemical feedstock pales in comparison to its global production.101 In this study, we focus on the economic and environmental trade-offs of producing benzene, toluene, and/or the xylenes from biomass via methanol, since it is a versatile product that can be converted from biomass via the synthesis gas route. We propose the f irst quantitative f ramework capable of systematically comparing several novel and commercial technologies to determine the optimal process topology for the production of aromatics from biomass. The process synthesis superstructure proposed in this paper constitutes a large-scale, mixed-integer, nonlinear, nonconvex optimization model that is solved to global optimality using a branch-and-bound global optimization framework. Simultaneous heat, power, and water integration is included within the framework to minimize the intake of freshwater into the refinery and convert waste heat into electricity.126−133 Each unit within the process superstructure is rigorously modeled to ensure

proper operation. Several novel, commercial, and/or competing technologies are modeled within the framework, including methanol-to-aromatics, toluene alkylation with methanol, selective toluene disproportionation, toluene disproportionation and transalkylation of toluene with C9+ aromatics, p-xylene separation via adsorptive separation or crystallization, isomerization of xylenes, dehydrocyclodimerization of liquefied petroleum gas, and separation via conventional distillation, to name a few. The optimal process topology, economic and environmental trade-offs, and life-cycle emissions analysis are determined and discussed for several case studies. The process synthesis framework for the BTA process will include (i) biomass gasification with/without recycle light gas, (ii) synthesis gas treatment, (iii) synthesis gas conversion via methanol synthesis, (iv) methanol conversion via methanolto-aromatics or methylation of toluene, (v) hydrocarbon upgrading, (vi) aromatization over a Ga/H-ZSM-5 based catalyst, and (vii) separation and production of additional high-value chemicals via an aromatics complex. The major products will be benzene, toluene, p-xylene, m-xylene, and/or o-xylene. Gasoline, LPG, and/ or electricity are allowable byproducts. A description of the major sections of the refinery is included in the following section.

2. BTA PROCESS SUPERSTRUCTURE: CONCEPTUAL DESIGN AND MATHEMATICAL MODELING The following subsections will describe the major components of the BTA refinery. In each of the figures illustrated throughout this section, input and output units are colored light green, units represented with a binary variable are dark blue, all other units are light blue, variable process streams are blue, and existing process streams are gray. The complete mathematical model of the process synthesis superstructure is included in Appendix B. 2.1. Biomass Handling and Gasification. Three different categories of biomass are investigated to determine the effect of biomass type on the overall profit of the refinery. The types investigated include agricultural residues (corn stover), perennial crops (switchgrass), and forest residues (hardwood). The refinery will input only one type of biomass to increase uniformity and decrease complexity in the preprocessing steps. Forest residues are delivered as woodchips to the refinery and screened to direct any sizes larger than 2 in. to a grinder before entering the gasifier. Agricultural residues and perennial crops are delivered as bales and grinded before being sent to the gasifier. The proximate and ultimate analyses for the biomass types are shown in Table 1.134 Table 1. Feedstock Proximate and Ultimate Analysis for Biomass Species134 proximate analysis (db, weight %) feed type

moist. (ar)

agricultural perennial forest

6.1 8.2 45

HHVc

LHVd

5.1 80.9 14 18101 4.6 79.2 16.2 18636 2.14 N/A N/A 19130 ultimate analysis (db, weight %)

16849 17360 17842

ash

VMa

FCb

heating values (kJ/kg)

feed type

C

H

N

Cl

S

O

agricultural perennial forest

46.8 46.9 50.19

5.74 5.85 5.9

0.66 0.58 0.32

0.266 0.501 0

0.11 0.11 0.03

41.4 41.5 41.42

VM = volatile matters. bFC = fixed carbon. cHHV = higher heating value. dLHV = lower heating value. a

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Figure 1. Biomass gasification flow sheet.

The biomass gasification flow sheet is shown in Figure 1. The biomass drier is modeled using a binary variable and is activated if the biomass moisture content is greater than 20 wt % in order to reduce the moisture down to this threshold. Otherwise, the biomass bypasses this unit. Flue gas generated within the refinery supplies the heat necessary for drying. The flue gas exits the drier at 110 °C and 1.05 bar and is passed through an air cyclone and baghouse filter so that any particulates may be removed.135 The biomass is lockhopped using compressed CO2 (10 wt %) and is transferred to the gasifier. The biomass gasifier operates at 30 bar and at a temperature of either 900 °C, 1000 °C, or 1100 °C. The input into the gasifier will either be solid biomass or a mixture of solid biomass and recycle light gases and is modeled using the stoichiometry-based model proposed by Baliban et al.128,136 The composition of the effluent is a function of the biomass gasifier temperature, oxidizer flow rate, and biomass composition. The gasifier effluent will contain a mixture of synthesis gas (H2, CO, CO2, H2O), C1 and C2 hydrocarbons, ash, tar, char, NH3, and acid gases such as H2S.128,136 The solid ash and char are separated from the vapor phase effluent using two biomass cyclones and are recycled back to the gasifier. This ensures that effectively 100% of the carbon present in the biomass feedstock is converted. The ash is ultimately output as slag. Steam input into the biomass gasifier will reform the C1−C2 hydrocarbons and the tar species, as well as gasify the feed. The heat necessary for the reforming reactions will be provided by high-purity oxygen, which additionally serves to facilitate the cracking of the tar species. Since the biomass gasifier operates at high temperatures, the syngas species will approach the water− gas-shift (WGS) equilibrium. Hydrocarbons, however, will be above their equilibrium values. Additional CO2 produced within the refinery can be recycled back to the biomass gasifiers. However, this will require generation of H2 from either pressureswing adsorption or electrolysis of water to aid in the conversion of CO2 via the reverse water−gas-shift reaction. The effluent exiting the biomass ash cyclones is sent to a catalytic tar cracker operating at 825 °C. The tar cracker is modeled after

one presented by the National Renewable Energy Laboratory.135 A circulating catalyst between the tar reformer and the catalyst regenerator supplies a portion of the heat necessary for tar conversion.137 Additional combustion gases are sent to the regenerator to provide the balance of heat for the endothermic tar reforming reactions.137 After exiting the tar cracker, the biomass syngas is directed to the syngas cleaning section. 2.2. Synthesis Gas Cleaning. The synthesis gas cleaning section for the BTA refinery is shown in Figure 2. The biomass synthesis gas can either be split to a dedicated sour water− gas-shift reactor or can bypass this unit. The sour water−gas-shift reactor operates at a pressure of 26 bar and at a temperature of either 300 °C, 400 °C, 500 °C, or 600 °C. This unit can either be used as a forward or reverse WGS reactor. Any CO2 generated within the refinery can be consumed via the reverse water− gas-shift reaction. Hydrogen produced within the refinery will help facilitate the consumption of CO2. Heat necessary for the reverse water−gas-shift reaction will be provided by combusting a portion of the synthesis gas species via oxygen input into the reactor. The forward water−gas-shift reaction provides a means of increasing the H2/CO ratio of the syngas at the expense of producing CO2. Since the forward water−gas-shift reaction is exothermic, low pressure steam will be generated. Reformed gases from the autothermal reformer are also allowed to enter the water−gas-shift reactor. The existence of the dedicated WGS reactor depends on the H2/(CO + CO2) ratio needed for syngas conversion, the amount of H2 that can be economically produced using the PSA, the economics surrounding H2 production from electrolysis, the investment costs associated with the WGS reactor, and the specified life cycle greenhouse gas emissions of the refinery. If the WGS reactor is included in the optimal topology, the syngas effluent will be cooled down to 185 °C and will be directed to the scrubbing section. Otherwise, the synthesis gas exiting the tar cracker is cooled down to 185 °C before being sent to the scrubbing section. The scrubbing section removes residual tar, particulates, and NH3 that is contained within the raw synthesis gas. The wastewater stream from this section is directed to the wastewater C

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Figure 2. Biomass syngas cleaning flow sheet.

Figure 3. Methanol synthesis and conversion flow sheet.

treatment section. The scrubbed synthesis gas is directed to a Rectisol unit that removes the acid gases from the synthesis gas. This unit is necessary within the refinery to prevent contamination in the downstream hydrocarbon production and upgrading sections. The effluent from the Rectisol unit contains a sulfur-rich stream that is directed to the Claus recovery system to convert 95% of the H2S and SO2 into solid sulfur.138 The tail gas from the Claus system is hydrogenated to H2S, thus effectively achieving 100% recovery of the sulfur in the refinery. The carbon dioxide from the acid gas recovery units can either be (i) compressed to 31 bar and recycled within the refinery, (ii) compressed to 150 bar for sequestration, or (iii) vented to the atmosphere. 2.3. Hydrocarbon Production & Upgrading. 2.3.1. Methanol Synthesis. The methanol synthesis reactor operates at 250 °C and 45 bar and converts the synthesis gas generated

within the refinery to methanol via the water−gas-shift reaction and the methanol synthesis reaction, shown in eqs 1 and 2, respectively. The process flow diagram is shown in Figure 3. CO + H 2O ↔ CO2 + H 2

(1)

CO + 2·H 2 ↔ CH3OH

(2)

The effluent from the methanol synthesis reactor is cooled and flashed to remove the majority (96 mol %) of the methanol. The vapor effluent is split so that the majority (96 mol %) is recycled back to the methanol synthesis reactor, while the rest (4 mol %) is used as fuel gas within the refinery. The liquid effluent is heated to 371 °C to form a vapor and expanded in a turbine to generate electricity. The turbine effluent is cooled down and sent to a degasser distillation column. The degasser D

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assumed to be formed, and it is assumed that the distribution between the m- and o-xylene isomers is 68 wt % meta to 32 wt % ortho.141 Wastewater produced within the process is directed to the wastewater treatment section, the aromatics produced are sent for further processing in the aromatics complex, and the light gases are recycled within the BTA refinery. 2.3.3. LPG-Aromatics Separation. The LPG-aromatics separation section inputs the raw hydrocarbon product produced in the methanol-to-aromatics reactor. Light gases are initially knocked out from the hydrocarbon mixture, which is then directed to a deethanizer distillation column that separates any residual light hydrocarbons. The bottoms from the deethanizer distillation column are directed to the stabilizer column to remove the C3/C4 gases from the heavier hydrocarbons. The C3/C4 gases are directed to a HF alkylation reactor which converts the isobutane and butene into isooctane.135 The effluent from the HF alkylation reactor is directed to the LPG/Alkylate splitter, which separates out the byproduct LPG. The bottoms from the stabilizer column is directed to a splitter column, which provides lean oil for use in the absorber column. The effluent from the splitter column is directed to the aromatics complex since it contains a mixture of heavier aliphatics and aromatics. The tops from the deethanizer column is directed to an absorber column which uses the lean oil to remove any C3+ species.135 The light gases are ultimately recycled back into the refinery, while the lean oil and absorbed hydrocarbons are refluxed back to the deethanizer column. The LPG-aromatics flow diagram is shown in Figure 4 and is based off one presented by the National Renewable Energy Laboratory.135 2.4. LPG Processing. The LPG processing section is shown in Figure 5. LPG generated within the refinery can either be output as byproduct, recycled back in the refinery and utilized, or can undergo dehydrocyclodimerization over a metal promoted H-ZSM-5 catalyst to produce high-value aromatics, as is the case in the Cyclar process.97,144 The Cyclar process utilizes a Ga/H-ZSM-5 catalyst to convert LPG into aromatics via dehydrogenation and subsequent aromatization.97,144 Developed by BP and UOP, the first commercial scale Cyclar process was commissioned by BP in 1990. Along with high-value aromatics, a valuable coproduct, hydrogen, is also produced. The decision of whether to include the Cyclar process within the BTA refinery depends on the total cost associated with this unit, as well as the relative selling prices of LPG and aromatics. The branch-and-bound global optimization framework will determine the optimal split fraction to send to the Cyclar process. The hydrogen and fuel gas are recycled back within the refinery, while the aromatics are directed to the aromatics complex. 2.5. Aromatics Complex. Several commercial and competing technologies are included within the aromatics complex to determine the optimal configuration for the upgrading and separation of high-value C6−C8 chemicals. The following subsections will describe these technologies. The process flow diagram for the aromatics complex is shown in Figure 6. 2.5.1. Sulfolane Unit. The UOP Sulfolane process recovers aromatics from reformate, pyrolysis gasoline, and coke-oven light oil using a combination of liquid−liquid extraction and extractive distillation. The solvent commonly used is tetrahydrothiophene 1,1-dioxide, or sulfolane.97,145 The process recovers 99.9% of the benzene, 99.8% of the toluene, and 99% of the C8 aromatics from the incoming stream. The MTA effluent will contain some paraffins that are separated and directed to the gasoline blender. Before being directed downstream for further fractionation,

removes the entrained gases from the methanol/water mixture. The light gases are recycled back to the refinery as fuel gas. The methanol/water mixture is split to either the methanolto-aromatics reactor or the GTC GT-TolAlk toluene methylation reactor. 2.3.2. Methanol Conversion. The methanol/water mixture exiting the methanol degasser can be processed in one of two options, the methanol-to-aromatics (MTA) reactor or the toluene methylation reactor. Prior to entering the MTA reactor, the degasser effluent is heated up to 425 °C. Several metal-promoted ZSM-5 catalysts, such as Zn/ZSM-5, Ga/ZSM-5, and Ag/ZSM-5, are known to increase overall selectivity of methanol to aromatics.112,114,116 The methanol enters the MTA reactor and is converted into an aromatic-rich hydrocarbon stream over a Ag-ZSM-5 catalyst. The overall reaction is shown in eq 3. The reactor operates at 425 °C, and the methanol is completely converted into 56 wt % water and 44 wt % hydrocarbons.116 The MTA reactor is modeled using an atom balance assuming that the hydrocarbon product is distributed according to Table 1 and Table 2 of Inoue et al. at 700 K.116 For convenience, the data Table 2. Product Distribution (in C%) for the Methanol to Aromatics Reaction Using a Ag/ZSM-5 Catalyst116 temperature (°C) product distribution in C% CH4 C2H6 C3H8 C4H10 C2H4 C3H6 C4H8 C5+ aliphatics C6H6 C7H8 e-C8H10 m,p-C8H10 o-C8H10 C9H12 aromatics other aromatics

425 1.3 0.4 3.5 4.0 8.1 6.2 2.5 1.5 1.6675 7.6125 0.5075 23.9975 7.105 25.955 5.655

from Inoue et al. is shown in Table 2.116 Other aromatics are modeled as C10H14 aromatic compounds. As eq 3 illustrates, 0.22 mol of H2 will be produced for every mole of methanol converted. The MTA effluent is upgraded and separated in the LPG-Aromatics Separation section, which is modeled after the scheme presented by the National Renewable Energy Laboratory and is described later in the text.135 The LPG-Aromatics separation section separates the light gases, LPG, and higher hydrocarbons; ultimately, a portion of the higher nonaromatic hydrocarbons are sent to the gasoline pool. CH3OH → CH1.56 + 0.22·H 2 + H 2O

(3)

Alternatively, a portion of the methanol may be split to a toluene methylation reactor that uses a proprietary zeolite catalyst developed by GTC Technology Corporation.139−142 The methanol is mixed with toluene produced within the refinery and is reacted over a ZSM-5 based catalyst to produce fuel gas, mixed xylenes, C9 aromatics, and process water that are separated downstream.139 The weight ratio of toluene to methanol input into the reactor is 4:1.89.139,142,143 Xylenes are produced with a selectivity of 85 wt % to p-xylene.139,142,143 No ethyl-benzene is E

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Figure 4. LPG-aromatics separation flow sheet. Reprinted from ref 107 with permission. Copyright 2016 Wiley.

Figure 5. LPG processing flow sheet.

the aromatic-rich effluent is clay treated to remove any impurities.100 The Sulfolane process was developed by UOP in the 1960s. The aromatic-rich stream from the UOP Sulfolane process is mixed with any other aromatic-rich streams within the refinery and directed to the benzene distillation column. 2.5.2. Benzene Distillation Column. As shown in Figure 6, the benzene distillation column inputs a stream of C6+ aromatics and separates high-purity (99.8 wt %) benzene in the distillate. The distillation column recovers 99.9% of the incoming benzene in the distillate, which is then cooled to 40 °C before being output as benzene product. The benzene depleted bottom is directed to the toluene distillation column. 2.5.3. Toluene Distillation Column. The toluene distillation column recovers high-purity (99.8 wt %) toluene in the distillate, which also contains any leftover benzene and a small portion of C8 aromatics. The toluene distillate can then be split among five different processing options. The toluene distillate can be (1) output as toluene product, (2) directed to the UOP Tatoray reactor, (3) directed to the UOP TAC9 reactor, (4) directed to the UOP PX-Plus XP process, or (5) directed to the GTC GT-TolAlk reactor. The optimal split fraction to each of these alternatives depends on the selling price of toluene, the investment and operating costs of the alternatives, and the composition of the products formed. This decision will be determined by the deterministic global optimization branch-and-bound framework.

The toluene depleted bottom is mixed with any other C8+-rich aromatic stream in the refinery and is then directed to either one of two xylene distillation columns, described below. 2.5.4. Xylene Column 1. The first xylene distillation column allows for the production of o-xylene within the BTA refinery. Since the boiling point of o-xylene is higher than those of the other C8 isomers, as shown in Table 3, separation of o-xylene from these compounds via distillation is possible.146 Xylene column 1 recovers p-xylene, m-xylene, and ethylbenzene (all with greater than 99.9% recovery) in the distillate stream. The distillate is directed to a splitter that determines the optimal ratio of C8 aromatics to be directed to the UOP Parex process. The bottom stream contains o-xylene and heavier aromatics that are directed to the ortho-distillation column. The production of o-xylene within the refinery must proceed through this route. 2.5.5. Ortho-Xylene Distillation Column. The bottom from xylene column 1 is directed to the o-xylene distillation column, which recovers high purity (99.5%) o-xylene in the distillate stream. The o-xylene distillation column recovers 99.9% of the incoming o-xylene. The bottom from the o-xylene distillation column contains mostly C9+ aromatics and is either split to the gasoline pool or to the C9 distillation column. 2.5.6. Xylene Column 2. Alternatively, the C8+-rich aromatic stream can be directed to xylene column 2, which recovers over 99.4% of the p-xylene, o-xylene, m-xylene, and ethylbenzene from the incoming feed in the distillate. The distillate is mixed with any F

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Figure 6. C6−C8 aromatics upgrading and separation flow sheet.

cooled down to 40 °C before being output as product. The p-xylene depleted effluent is split to either the UOP MX Sorbex process to selectively recover m-xylene or to the UOP Isomer process, both of which are described below. 2.5.8. UOP MX Sorbex. m-Xylene is recovered from the other C8 isomers using the UOP MX Sorbex process, a continuous adsorptive separation system which recovers 95% of the incoming m-xylene.149 The m-xylene effluent has a mass purity of at least 99.5% and is cooled down to 40 °C before being output as product. The m-xylene depleted stream is sent to the UOP Isomar to reestablish an equilibrium mixture of xylene isomers. 2.5.9. UOP Isomar. The UOP Isomar process uses an I-400 EB isomerization catalyst to reestablish an equilibrium mixture of xylene isomers and convert any ethylbenzene within the feed to mixed xylenes with 30% wt per pass conversion.97,100 An alternative catalyst, the I-300 EB dealkylation catalyst, exists and converts ethylbenzene to benzene with greater than 70% wt per pass conversion.97,100 However, it is not modeled within the BTA process superstructure. The UOP Isomar allows for the exclusive production of either p-xylene, m-xylene, and o-xylene, an important consideration that will be depicted within the case studies later in the paper. The effluent from the UOP Isomar process is sent to a deheptanizer distillation column that purges the C7 light ends. The C8-rich aromatic mixture is then mixed with the bottoms from the toluene distillation column and directed to one of the two xylene distillation columns.

Table 3. Boiling Points (°C) of C8 Aromatics146 C8 isomer

boiling point

ethylbenzene m-xylene o-xylene p-xylene

136.19 139.10 144.41 138.35

other C8-rich aromatic stream and is sent to a splitter, which determines the optimal ratio of C8 aromatics to be directed to the UOP Parex process. The bottom from this distillation column is split to either the C9 distillation column or directed to the gasoline blender. Xylene column 1 and xylene column 2 are modeled with binary variables such that only one of the columns is allowed to exist within the refinery. 2.5.7. UOP Parex. p-Xylene may be selectively recovered from a mixture of C8 isomers using an adsorptive separation system (also known as a simulated moving bed technology), as is the case in the UOP Parex process.147,148 Since the boiling point of p-xylene is close to the boiling points of the other isomers, separation of p-xylene using conventional distillation is not practical and typically occurs through either fractional crystallization or adsorption. The Parex process consists of two adsorption chambers with 12 beds each and utilizes a solid zeolitic adsorbent to selectively separate para-xylene.97 The process operates at a temperature of 180 °C and 1.2 bar and recovers 99% of the incoming paraxylene with a mass purity of at least 99.5%. The p-xylene effluent is G

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other C8 isomers. The first alternative, the UOP Parex process, was described above. Before the introduction of the Parex system, the only method of separating p-xylene from the other C8 aromatic isomers was via fractional crystallization. However, the effective separation of p-xylene using crystallization is impaired because of a eutectic composition limit that only allows for 65% per pass recovery of p-xylene.97 This is mediated, however, when the feed into the crystallization system contains over 70% p-xylene.156 Because of this, the crystallization process is coupled with the UOP PX-Plus process. The single-stage crystallization process recovers high-purity p-xylene (99.9%). The effluent from the PX-Plus is directed to an upgrading process that separates out benzene and heavier aromatics. Light gases exiting the PX-Plus are recycled in the refinery, and the C9 aromatics are split and either output as gasoline or directed to the UOP Tatoray or UOP TAC9 processes. The Badger/Niro crystallization process outputs high-purity para-xylene, while the mother liquor (a mixture of C8 aromatic isomers) is directed to the UOP Parex process. The effluent exiting the PX-Plus reactor is composed of 7.5 wt % light gases, 43.4 wt % benzene, 35.6 wt % p-xylene, 11.4 wt % mixed xylenes, and 2.1 wt % C9 aromatics.156 The mixed xylenes are decomposed according to Table 3 of an excellent review by Tsai et al. and the information in the PX-Plus packet provided by UOP.102,156 Ernst also provides a typical product distribution for the PX-Plus reactor.142 2.6. Light Gas Handling. Light gas that is generated within the BTA refinery can be recycled back within the process in order to increase overall conversion, produce heat for process units, or generate electricity. An internal or external gas loop configuration is designed in order to handle the light gases produced within the BTA refinery. The internal gas loop configuration will recycle light gases produced via methanol synthesis back to the reactor to increase the overall conversion. A portion of these light gases must be purged to prevent buildup of inert species within the refinery. The light gases purged from methanol synthesis, along with any other light gases generated within the refinery, will be handled via the external gas loop configuration. The external gas loop configuration will direct the light gases to either the fuel combustor to provide heat for process units, the gas turbine to generate electricity, or an autothermal reformer to reform the light gases into syngas components that can be recycled back in the refinery. The effluent from the fuel combustor and the gas turbine can either be vented or cooled down to 35 °C, passed through a water knockout unit, and sent to a CO2 capture unit. The reformed gases from the autothermal reformer can be recycled to several process units within the BTA refinery, including the water−gas-shift reactor and the biomass gasifier, among others. The external gas loop configuration is shown in Figure 7. 2.7. Hydrogen/Oxygen Production. Several process units within the BTA superstructure require hydrogen and oxygen. In order to satisfy this requirement, high purity hydrogen will be provided via electrolysis of water or pressure-swing adsorption. An air separation unit or an electrolyzer will provide the necessary oxygen. The process flow diagram is shown in Appendix A (Supporting Information). 2.8. Wastewater Treatment. A comprehensive wastewater treatment section is included within the BTA superstructure to treat process wastewater.129 The process wastewater treatment section includes a biological digestor and sour stripper that inputs sour water and acid-rich wastewater from process units within the refinery. The utility wastewater treatment section outputs process water for electrolysis and steam to process units.

2.5.10. C9 Distillation Column. The BTA process superstructure currently contains three processing options for the heavy aromatics produced within the refinery. The heavy aromatics exiting the bottom of the o-xylene distillation column and xylene column 2 can either be split to the C9 distillation column or output as gasoline. The distillate from the C9 distillation column contains 99% of the incoming C9 aromatics and a small portion of the incoming C10 aromatics. The distillate is split between the UOP Tatoray process and the UOP TAC9 process, which are described in the subsequent subsections. The bottoms product, which predominantly contains a C10+-rich aromatic stream, is directed to the gasoline blender to be output as gasoline. 2.5.11. UOP Tatoray. Any toluene produced within the refinery can be reacted with C9 aromatics via transalkylation (eq 5) within the Tatoray reactor to produce mixed xylenes. Additionally, the disproportionation of toluene (shown in eq 4) to produce mixed xylenes and benzene also takes place.97,150 The UOP Tatoray process utilizes a fixed bed reactor that operates at a temperature of 430 °C and 35 bar.97,151 A 50−50 wt % ratio of toluene and C9 aromatics is input along with a small amount of makeup hydrogen. The UOP Tatoray process is modeled using an atom balance and the information provided in Table 1 of U.S. Patent 7,109,389 B2.151 The effluent from the Tatoray process is mixed with the effluent from the UOP TAC9 process and the UOP Sulfolane process and is sent to be fractionated, as shown in Figure 6. 2·C7H8 ↔ C6H6 + C8H10

(4)

C7H8 + C9H12 ↔ 2·C8H10

(5)

2.5.12. UOP TAC9. Alternatively, the distillate from the C9 distillation column can be split to the UOP TAC9 process, where the conversion of heavy aromatics into xylenes via transalkylation takes place in a fixed bed reactor. A mixture of C9−C10 aromatics can be input into the TAC9 process to be converted into mixed xylenes.152,153 The maximum amount of C9+ aromatics that the TAC9 process is able to handle is 70%;154 therefore, a 2.5:1 molar ratio of C9 aromatics and toluene is input. The reactor operates at a temperature of 300 °C, a pressure of 1.2 bar, and the per-pass conversion of toluene and C9 aromatics is roughly 40% and 45%, respectively.155 The composition of the TAC9 effluent is derived from Table 3 of a study by Krejči ́ et al.155 and is modeled using an atomic balance. The aromatic effluent exiting the reactor is mixed with the effluent from the UOP Sulfolane process and the UOP Tatoray process and is directed to the fractionation columns. The global optimization branch-and-bound framework will determine whether the UOP Tatoray process, the UOP TAC9 process, or both will exist within the BTA refinery. This decision depends on the relative amounts of C9+ aromatics and toluene exiting the hydrocarbon production section, the investment and operating costs of these processes, and their respective effluent compositions. 2.5.13. UOP PX-Plus XP Process. The UOP PX-Plus XP is the final alternative for the conversion of toluene within the biomass to aromatics process synthesis superstructure. The UOP PX-Plus XP process is a combination of the PX-Plus process, which selectively disproportionates toluene into p-xylene, and the Badger/Niro p-Xylene crystallization process, which recovers p-xylene from a stream of C8 isomers.142,156 The selectivity to the p-xylene isomer is close to 90%.142 Additionally, a small amount of H2 is input into the process. The PX-Plus process also produces a benzene coproduct. The Badger/Niro p-Xylene crystallization process represents the second alternative for the separation of p-xylene from the H

DOI: 10.1021/acs.energyfuels.6b00619 Energy Fuels XXXX, XXX, XXX−XXX

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Energy & Fuels

Figure 7. Light gas recycle flow sheet.

The process flow diagrams are shown in Appendix A (Supporting Information). 2.9. Unit Costs. The total direct costs (TDC) of each unit in the BTA refinery are rigorously calculated using eq 6: TDC = (1 + BOP)Co(

Sr sf ) So

4.5% of the total plant costs and the refinery is assumed to operate for 330 days/yr (CAP). These values are used to convert the total cost of a unit into its levelized unit cost, (costU), as shown in eq 8: cos t Uu =

(6)

(8)

where Prod levelizes the cost based on the total energy of products produced. 2.10. Utility Requirements. Several units within the BTA refinery have utility requirements that are already known, and thus are modeled with known energy balances and requirements. The utility requirements for these units are shown in Appendix A.1 (Supporting Information). The remaining units (e.g., heat exchangers, compressors, and others) are calculated using the conservation of energy and total heat balance equations shown in Appendix B (Supporting Information). The process synthesis framework determines the total minimum utility cost for the process. 2.11. Objective Function. Equation 9 illustrates the objective function for the MINLP model described in the previous sections:

where BOP represents the balance of plant (site preparation, civil works, etc.) percentage (20% of the total installed cost), Co is the base component cost, So is the base component size (capacity), Sr is the actual component size (capacity), and sf is the scaling factor. In order to accurately estimate the total cost of each unit, several literature sources16,32,97,135,138,139,152,156−159 are used to the determine the parameters used in eq 6. The Chemical Engineering Plant Cost Index is used to convert all costs in 2015 dollars.160 The cost parameters are shown in Table 4. Engineering, contingency, startup, royalties, fees, and spare parts are included within the indirect costs of the refinery.32 The indirect costs are assumed to be 32% of the total direct costs in the BTA refinery. The total plant costs (TPC), also referred to as the fixed capital investment (FCI), are calculated as the sum of the total direct costs and the total indirect costs. To levelize the total plant costs to a yearly cost, the levelized capital charge rate (LCCR) and the interest during construction factor (IDCF) are used in eq 7 to calculate the capital charges: CC = LCCR· IDCF· TPC

⎛ LCCR· IDCF OM ⎞ ⎛ TPCu ⎞ ⎜ ⎟·⎜ ⎟ + ⎝ CAP 365 ⎠ ⎝ Prod ⎠

MIN

∑

cos t f + cos tEl + cos tSeq +

f ∈ Feed

−

(7)

∑ p ∈ Products

The values of LCCR (14.38%/yr) and IDCF (7.16%/yr) are taken from Kreutz et al.32 The overall multiplier, 15.41%/yr, converts the total plant costs into total yearly capital charges. The operating and maintenance (OM) costs are assumed to be

∑

cos tUu

u ∈ UInv

cos t p (9)

The summation represents the total negative profit and includes contributions from the feedstock costs, costf, the electricity costs/profit, costEl, the CO2 sequestration cost, costSeq, the I

DOI: 10.1021/acs.energyfuels.6b00619 Energy Fuels XXXX, XXX, XXX−XXX

Article

Energy & Fuels Table 4. BTA Refinery Upgrading Unit Reference Capacities, Costs (2015 $), and Scaling Factors description

Co (MM$)

So

biomass handling (forest) biomass handling (nonforest) biomass gasification

$4.55 $14.11 $54.04

17.90 17.90 17.90

water-gas-shift unit rectisol unit

$3.67 $31.42

150.00 2.51

methanol synthesis and degasser methanol-to-aromatics unit GTC toluene alkylation

$8.05 $8.67 $42.42

35.65 10.63 13.89

deethanizer absorber column stabilizer column splitter column HF alkylation unit LPG/alkylate splitter autothermal reformer

$0.57 $0.89 $1.01 $0.99 $8.79 $1.04 $21.50

5.13 0.96 4.57 3.96 0.61 0.61 12.20

UOP sulfolane UOP isomar UOP Tatoray UOP Parex UOP MX Sorbex benzene distillation toluene distillation xylene1 column xylene2 column ortho-xylene distillation C9 distillation PXPLUS XP UOP TAC9 cyclar process

$21.15 $45.91 $22.25 $153.54 $96.54 $0.56 $1.16 $5.00 $2.19 $0.99 $0.93 $65.38 $16.06 $112.67

15.14 1555.50 24.82 106.20 5.58 10.83 10.04 12.11 11.72 12.14 15.31 19.38 13.19 16.42

pressure-swing adsorption air separation unit air compressor electrolyzer

$7.79 $243.50 $5.90 $0.49

0.29 145.00 10.00 1.00

gas turbine steam turbine

$79.84 $64.87

266.00 136.00

sour stripper biological digestor reverse osmosis cooling tower

$3.91 $4.65 $0.31 $61.04

11.52 115.74 4.63 1.75

SMax

units

scale basis

Biomass Conversion 33.30 kg/s as received biomass 33.30 kg/s as received biomass 33.30 kg/s dry biomass Synthesis Gas Handling/Clean-Up 250.00 kg/s feed 8.78 kmol/s feed Hydrocarbon Production kg/s feed 42.50 kg/s feed 200.00 kg/s toluene feed Hydrocarbon Upgrading kg/s feed kg/s feed kg/s feed kg/s feed kg/s feed kg/s feed 35.00 kg/s feed Aromatics Production Complex 100.00 kg/s feed 200.00 kg/s feed 200.00 kg/s feed 200.00 kg/s feed 200.00 kg/s feed 200.00 kg/s feed 200.00 kg/s feed 200.00 kg/s feed 200.00 kg/s feed 200.00 kg/s feed 200.00 kg/s feed 200.00 kg/s toluene feed 200.00 kg/s feed 200.00 kg/s feed Hydrogen/Oxygen Production kmol/s purge gas 41.70 kg/s O2 feed 30.00 MW electricity MW electricity Heat and Power Integration 334.00 MW electricity 500.00 MW electricity Wastewater Treatment kg/s feed kg/s feed kg/s feed gW heating deficit

levelized investment cost, costU, and the profits from selling the products, costp. The feedstock costs include contributions from biomass and freshwater, while the product profits include contributions from benzene, toluene, p-xylene, m-xylene, o-xylene, gasoline, and LPG. Each term is normalized with respect to the total energy of products (aromatics/liquid fuels) produced. The objective function is minimized, which is equivalent to maximizing the total profit of the BTA refineries. 2.12. Deterministic Global Optimization. The process synthesis superstructure described in the preceding sections is modeled mathematically as a large-scale, nonconvex, mixedinteger, nonlinear optimization (MINLP) model. The resulting MINLP model contains 19 303 continuous variables, 28 binary

sf

ref

0.77 0.77 0.77

Kreutz et al., 200832 Kreutz et al., 200832 Larson et al., 200916

0.67 0.63

NREL, 2011135 Kreutz et al., 200832

0.65 0.65 0.67

NREL, 2011135 NREL, 2011135 GTC, 2014139

0.68 0.68 0.68 0.68 0.65 0.68 0.67

NREL, 2011;135 Mobil R&D, 1978157 NREL, 2011;135 Mobil R&D, 1978157 NREL, 2011;135 Mobil R&D, 1978157 NREL, 2011;135Mobil R&D, 1978157 NREL, 2011;135 Mobil R&D, 1978157 NREL, 2011;135 Mobil R&D, 1978157 NETL, 2013158

0.67 0.67 0.67 0.67 0.67 0.52 0.72 0.49 0.77 0.66 0.53 0.67 0.67 0.67

Meyers, 200497 Meyers, 200497 Meyers, 200497 Meyers, 200497 Meyers, 200497 Aspen Plus PEA Aspen Plus PEA Aspen Plus PEA Aspen Plus PEA Aspen Plus PEA Aspen Plus PEA UOP, 2006156 Hydrocarbon Processing, 2010152 Meyers, 200497

0.65 0.50 0.67 0.90

Larson et al., 200916 NETL, 2013158 Larson et al., 200916 Larson et al., 200916

0.75 0.67

Larson et al., 200916 Larson et al., 200916

0.53 0.71 0.85 0.65

NETL, 2010138 Balmer, 1994159 Balmer, 1994159 NETL, 2010138

variables, 23 463 constraints, and 528 nonconvex terms. The nonconvex terms consist of 468 bilinear terms, 1 trilinear term, 3 quadrilinear terms, and 56 concave power functions. In order to solve the model, a novel branch-and-bound161 algorithm is used and refined to generate satisfactory initial starting points that are used to find upper bound solutions. Any nonlinearity within the model is underestimated with its linear relaxation, and the resulting mixed-integer linear optimization (MILP) model is solved using CPLEX162 to generate the lower bound of the overall MINLP model. Bilinear terms are underestimated using piece-wise McCormick underestimators, where the number of pieces used depends logarithmically on the number of binary variables. Concave power functions used to represent the J

DOI: 10.1021/acs.energyfuels.6b00619 Energy Fuels XXXX, XXX, XXX−XXX

Article

Energy & Fuels investment cost of each individual unit are underestimated using piece-wise linear relaxations where each piece is represented using a binary variable. The solution of the lower bound (MILP model) generates several initial starting points. The binary variables are fixed at the starting points, and the resulting nonlinear optimization model (NLP) is solved using CONOPT163 to generate upper bound solutions. An optimality-based bounds tightening (OBBT) routine is formulated at the root node to tighten all variables associated with the total molar flow rates and nonlinear investment cost. After the completion of the OBBT route, a feasibility-based bounds tightening (FBBT) routine is started to additionally tighten all species molar flow rates. Each node in the branch-and-bound tree branches to form two children nodes, and an OBBT routine is again completed on selected variables before the start of each node. The upper bound is updated if the solution of the NLP is lower than the current upper bound. Nodes with a lower bound greater than the current upper bound are eliminated. For a detailed explanation of deterministic global optimization theory and algorithm, the reader is directed to textbooks by Floudas.164,165

Table 5. Cost Parameters for the BTA Refinery

a

commodity

nominal prices

agricultural (corn stover) perennial (switchgrass) forest (hardwood) freshwater electricity CO2 TS&Ma propane gasoline p-xylene m-xylene o-xylene benzene toluene

$120/dry metric ton $100/dry metric ton $70/dry metric ton $0.50/metric ton $0.07/kWhr $5/metric ton $0.99/gal $2.86/gal $1.602/kg $1.405/kg $1.367/kg $1.431/kg $1.294/kg

TS&M - transportation, storage, and monitoring.

The minimum number of heat exchanger matches is calculated using the above information as previously described.127−129,164,168 A superstructure approach is then utilized to determine the minimum investment cost of the heat exchanger network.127,129,164,168 The minimum cost of the heat exchanger network is then added to the overall BTA refinery investment cost. 3.1. Optimal Process Topologies. Table 6 shows the optimal process topologies selected for the biomass to aromatics refineries using nominal prices. Table 6 illustrates (i) the type of gasifier selected, (ii) the operating temperature of the biomass gasifier, (iii) the existence and operating temperature of the forward/reverse water−gas-shift reactor, (iv) the selection of the methanol conversion units, (v) the existence of the ortho-xylene distillation column, (vi) the existence of the UOP Parex process, (vii) the existence of the UOP Tatoray process, (viii) the existence of the UOP TAC 9 process, (ix) the existence of the UOP PXPLUS XP process, (x) the existence of the UOP Cyclar process, (xi) the selection of a CO2 sequestration system, and (xii) the existence of a gas turbine. As explained earlier, the input into the biomass gasifier can either be solid biomass or a mixture of solid biomass and recycle gases. Additionally, the biomass gasifier can operate at a temperature of either 900 °C, 1000 °C, or 1100 °C. In each of the corn stover and switchgrass BTA refineries, the input included a mixture of solid biomass and recycle gases. The only refineries that input solid biomass into the gasifier are the hardwood M-200 and O-500 refineries. In each of the refineries investigated, the optimal operating temperature of the gasifier was selected to be 900 °C. At lower operating temperatures, the gasifier requires less oxygen for combustion and produces less waste heat for steam generation. Additionally, the equilibrium constant of the forward water−gas-shift reaction is higher at lower temperatures, thus resulting in less favorable conditions for CO2 consumption. Since the synthesis gas produced from biomass has a H2/CO ratio that is less than ideal for methanol production, it needs to pass through a forward water−gas-shift reactor operating at 300 °C. All of the methanol exiting the methanol synthesis reactor was directed to the MTA (methanol-to-aromatics) reactor. The decision to exclude the GTC-TolAlk process was dependent on the operating and capital costs of the process and the composition of the effluent leaving the reactor. The types of aromatics to produce and the relative ratios of these chemicals is an important topological decision. In all of the case studies investigated, benzene was always produced as it is a major product in the methanol-to-aromatics reaction and currently no benzene conversion reactors exist within the superstructure. In all of the unrestricted case studies, the only C8 aromatic produced was para-xylene. The CS-U-100 and CS-U-1000 case studies also produced a toluene product. In the o-xylene case studies, only the CS-O-1000 and CS-O-2000 produced a toluene product. All of the m-xylene refineries inputting corn stover and switchgrass produced toluene. The only hardwood m-xylene refineries that produced toluene were the M-500 and M-1000.

3. COMPUTATIONAL STUDIES The process synthesis model described above (see Appendix B (Supporting Information) for its full description) is used to examine 45 distinct case studies across three sets of feedstocks: corn stover, switchgrass, and hardwood biomass. Three sets of aromatics products are considered that (a) place no restriction on the type of xylenes produced, (b) limit xylenes output to only the o-xylene isomer, and (c) limit xylenes output to only the m-xylene isomer. These case studies reflect the capability of our superstructure-based approach to produce any ratio of the p-, o-, and m-xylene isomers. The effect of refinery capacity is investigated by constraining the total output from the refinery to be 100, 200, 500, 1000, or 2000 t per day of xylene equivalent (based on its lower heating value) chemicals. The refinery capacities were selected such that a feasible amount of biomass could be economically transported to the plant gate every day. Allowable byproducts include gasoline, LPG, and electricity. A 75% minimum mass threshold for aromatics output was imposed in each of the case studies to enforce the stand-alone chemicals concept. The case studies will be denoted as N−C, where N represents the aromatics composition in the unrestricted (U), o-xylene (O), and m-xylene (M) refineries, and C represents the capacity in metric tons per day. Throughout the text, descriptions such as F−N−C may be used to differentiate between the corn stover (CS), switchgrass (SG), and hardwood (HW) refineries. The MINLP model includes a constraint that imposes at least a 50% reduction in life cycle greenhouse gas emissions from the biomass-based aromatics facility. This reduction is based on typical petroleum-based refineries (91.6 kg CO2eq/GJLHV),166 a high-value chemicals (HVC) production plant (1.7 kg CO2eq/kg of HVC),167 and a natural gas combined cycle plant (101.3 kg CO2eq/GJ) that produces electricity.138 If electricity is input (output) into the refinery, the GHG emissions are added (subtracted) to the life-cycle emissions of the BTA refinery. Nominal cost parameters used in the process synthesis model are illustrated in Table 5. Feedstock costs include costs associated with delivery to the plant gate. Product costs do not include costs associated with transportation of the products from the plant gate. The costs associated with the capture and compression of CO2 are included in the investment costs of the BTA refinery, shown in Table 4. The costs shown in Table 5 include the transportation, storage, and monitoring costs of CO2. The global optimization algorithm was allowed to run for 100 CPU hours or until all nodes in the branch-and-bound tree were processed. After completion of the global optimization branch-and-bound algorithm, the optimal BTA process topology will yield (i) the operating conditions of the heat engines, (ii) the amount of electricity produced by the heat engines, (iii) the amount of cooling water necessary, and (iv) the locations of all pinch points in the heat exchanger network. K

DOI: 10.1021/acs.energyfuels.6b00619 Energy Fuels XXXX, XXX, XXX−XXX

L

Biomass Conv. BGS Temp. WGS/RGS Temp. MTA Usage GTC-TolAlk usage ortho distillation UOP MX Sorbex UOP Parex UOP Tatoray

biomass conv BGS temp WGS/RGS temp MTA usage GTC-TolAlk usage ortho distillation UOP MX Sorbex UOP Parex UOP Tatoray UOP Tac9 UOP PXPLUS XP UOP Cyclar CO2SEQ Usage GT Usage

biomass conv. BGS Temp. WGS/RGS Temp. MTA usage GTC-tolAlk Usage ortho distillation UOP MX Sorbex UOP Parex UOP Tatoray UOP Tac9 UOP PXPLUS XP UOP Cyclar CO2SEQ usage GT usage

Y Y

Y Y

U-200

U-100

S/V 900.00 300.00 Y

Y

Y

S/V 900.00 300.00 Y

Y Y Y

U-200

U-100

Y Y Y

Y

Y

S/V 900.00 300.00 Y

Y Y Y

Y Y Y

S/V 900.00 300.00 Y

S/V 900.00 300.00 Y

U-200

S/V 900.00 300.00 Y

U-100

Y Y

S/V 900.00 300.00 Y

U-500

Y

Y Y Y

S/V 900.00 300.00 Y

U-500

Y

Y Y Y

S/V 900.00 300.00 Y

U-500

Y Y

S/V 900.00 300.00 Y

U-1000

Y

Y Y Y

S/V 900.00 300.00 Y

U-1000

Y

Y Y Y

S/V 900.00 300.00 Y

U-1000

Y Y

S/V 900.00 300.00 Y

U-2000

Y

Y Y Y

S/V 900.00 300.00 Y

U-2000

Y

Y Y Y

S/V 900.00 300.00 Y

U-2000

Y

Y

Y

Y

S/V 900.00 300.00 Y

O-200

Switchgrass

Y

S/V 900.00 300.00 Y

O-100

Y

Y Y

Y

Y

Y Y

S/V 900.00 300.00 Y

O-200

Hardwood

Y

S/V 900.00 300.00 Y

O-100

Y

Y Y

Y

Y

Y Y

S/V 900.00 300.00 Y

O-200

Corn Stover S/V 900.00 300.00 Y

O-100

Y

Y

S/V 900.00 300.00 Y

O-500

Y

Y Y

Y

S 900.00 300.00 Y

O-500

Y

Y Y

Y

S/V 900.00 300.00 Y

O-500

Y

Y

S/V 900.00 300.00 Y

O-1000

Y

Y Y

Y

S/V 900.00 300.00 Y

O-1000

Y

Y Y

Y

S/V 900.00 300.00 Y

O-1000

Table 6. Topological Information for the Optimal Solutions for Each of the Case Studies Is Displayed Belowa

Y

Y

S/V 900.00 300.00 Y

O-2000

Y

Y Y

Y

S/V 900.00 300.00 Y

O-2000

Y

Y Y

Y

S/V 900.00 300.00 Y

O-2000

M-100

Y

Y

S/V 900.00 300.00 Y

M-100

Y

Y Y

Y

S/V 900.00 300.00 Y

M-100

Y

Y Y

Y

S/V 900.00 300.00 Y

M-200

Y

Y

S/V 900.00 300.00 Y

M-200

Y

Y Y

Y

S 900.00 300.00 Y

M-200

Y

Y Y

Y

S/V 900.00 300.00 Y

M-500

Y

Y

S/V 900.00 300.00 Y

M-500

Y

Y Y

Y

S/V 900.00 300.00 Y

M-500

Y

Y Y

Y

S/V 900.00 300.00 Y

Y

Y

S/V 900.00 300.00 Y

M-1000

Y

Y Y

Y

S/V 900.00 300.00 Y

M-1000

Y

Y Y

Y

S/V 900.00 300.00 Y

M-1000

Y

Y

S/V 900.00 300.00 Y

M-2000

Y

Y Y

Y

S/V 900.00 300.00 Y

M-2000

Y

Y Y

Y

S/V 900.00 300.00 Y

M-2000

Energy & Fuels Article

DOI: 10.1021/acs.energyfuels.6b00619 Energy Fuels XXXX, XXX, XXX−XXX

Article

Biomass conversion (Biomass Conv.) can proceed through either gasification of the solid biomass (S) or a mixture of solid biomass and recycle gases (S/V). Temperature of the biomass gasifier (BGS temp; °C), along with the operating temperature of the forward/reverse water−gas-shift reactor (WGS/RGS temp; °C), is shown. The conversion of methanol using either the methanol-toaromatics (MTA) process or toluene methylation technologies (GTC-TolAlk) will be denoted using yes (Y) or no (−). Production of o-xylene using an o-xylene distillation column or production of mxylene using the UOP MX Sorbex reactor will be denoted using yes (Y) or no (−). Separation of p-xylene using the UOP Parex process will be denoted using yes (Y) or no (−). The existence of the UOP Tatoray, the UOP TAC9, or the UOP PX-Plus XP processes will be denoted using yes (Y) or no (−). The existence of the Cyclar unit to convert LPG into additional aromatics will be denoted using yes (Y) or no (−). Likewise, the utilization of a gas turbine (GT) or a CO2 sequestration system (CO2SEQ) is denoted using yes (Y) or no (−).

Toluene and heavy aromatics produced within the refinery were always split to both the UOP Tatoray and UOP TAC9 reactors. The UOP Tatoray reactor produces mixed xylenes and benzene, while the UOP TAC9 process produces xylenes. The UOP PX-Plus XP process was never selected. The LPG produced within the MTA reactor was always sent to the Cyclar process to produce additional aromatics, despite the added investment cost associated with this unit. All of the aromatics refineries were able to meet the environmental constraint (at least a 50% reduction in GHG emissions) imposed within the model without the additional investment and operating costs associated with a CO2 sequestration system. This is to be expected, since it is well-known that the utilization of biomass has significant environmental benefits. Finally, none of the case studies utilized a gas turbine. 3.1.1. Detailed Discussion on the Optimal Topology for the HW-U2000 Refinery. The process flow diagram for the HW-U-2000 case study is shown in Figure 8 as an illustrative example. Figure 9 shows the process flow diagram of the HW-U-2000 aromatics complex with corresponding stream flow rates shown in Table 7 for several relevant units. In both figures, the key topological decisions are shown. Note that several process units such as heat exchangers, compressors, and turbines are not shown. All of the methanol exiting the methanol synthesis reactor is split to the MTA reactor. The input into the toluene alkylation process requires toluene and methanol in a weight ratio of approximately 4:1.89. The global optimization branch-and-bound framework determined that it is more profitable to split the toluene (produced from the MTA reactor and processed in the aromatics complex) to the UOP Tatoray process and the UOP TAC9 process rather than include an additional expensive hydrocarbon production unit. Additionally, the selection of the UOP Tatoray process and UOP TAC9 process in the optimal topology allows for the conversion of heavy aromatics into more valuable mixed xylenes and benzene. The UOP PXPLUS-XP process is not selected in the optimal topology because of the input into the process and required investment cost. Although the PXPLUS process also produces mixed xylenes and benzene (43.4 wt % benzene, 35.6% p-xylene, 11.4 wt % mixed xylenes), it does not allow for the conversion of C9 aromatics within the refinery, which would then have to be directed to the gasoline pool if all toluene were split to the PXPLUS process. Additionally, the UOP Tatoray and the UOP TAC9 processes are both less expensive than the UOP PXPLUS-XP process. Even if the UOP PXPLUS-XP process were selected in the optimal topology, the UOP Parex process would still be required to separate p-xylene because the C8 aromatic streams within the refinery do not contain a large amount of p-xylene (i.e., > 70%). As Figure 9 and Table 7 show, 72.2% of the toluene distillate is split (from SPTOL to MXTATORAY) to the UOP Tatoray process, while the balance (27.8% - from SPTOL to MXTAC9) is directed to the UOP TAC9 process. 44.4% of the C9 distillate is split (from SPC9A to MXTATORAY) to the UOP Tatoray process, while the balance (55.6% - from SPC9A to MXTAC9) is directed to the UOP TAC9 process. The effluents from the UOP Tatoray and UOP TAC9 processes are mixed with the other aromatic-rich effluents from the refinery and directed to the benzene distillation column. Note that the H2 balance exiting the UOP Tatoray process (Table 7) is used to close the atom balances around the process and is not recycled back within the process.151 Also note that the values of the output streams in Figure 9 and Table 7 correspond exactly to the material balances shown in Table 11. The gasoline blend produced within the refinery contains mostly paraffins and heavy aromatics, along with some benzene, toluene, and xylenes. To meet all physical property specifications of gasoline, the aromatic-rich gasoline from the BTA refinery must be blended.135,169,170 The energy content for gasoline, which is used in Table 12, is based on a generic lower heating value of 5505.2 GJ per thousand barrels and a density of 747 kg/m3. 3.2. Overall Profit of Aromatics Production. The overall profit for each of the biomass to aromatics refineries is shown in Table 8. The overall profit is presented as the difference between the costs associated with the production of the aromatics/byproducts and the revenue from selling these products. Thus, the more negative the value

a

Y Y Y Y Y Y Y Y Y Y Y Y Y Y

Y

Y Y

M-1000 M-500

Y Y

M-200 M-100

Y Y

O-2000 O-1000

Y Y

O-500 O-200

Y Y Y Y Y Y Y

Switchgrass

O-100 U-2000 U-1000 U-500 U-200 U-100

Table 6. continued

UOP Tac9 UOP PXPLUS XP UOP Cyclar CO2SEQ Usage GT Usage

M-2000

Energy & Fuels

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DOI: 10.1021/acs.energyfuels.6b00619 Energy Fuels XXXX, XXX, XXX−XXX

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Energy & Fuels

Figure 8. Process flow diagram for the HW-U-2000 refinery.

Figure 9. Aromatics complex process flow diagram with stream information for the HW-U-2000 refinery. N

DOI: 10.1021/acs.energyfuels.6b00619 Energy Fuels XXXX, XXX, XXX−XXX

13

0 0 111.81203 233.83058 114.11079 9.95407 0.01076 0 0 0 0

C6H6 C7H8 o-C8H10 m-C8H10 p-C8H10 e-C8H10 C9H12 C10H14 C11H16 H2 H2 balance

C6H6 C7H8 o-C8H10 m-C8H10 p-C8H10 e-C8H10 C9H12 C10H14 C11H16

component

1

8.68695 27.69430 18.93431 42.96353 20.99561 1.32341 60.87253 12.18034 0

component

2

0 0 0.74433 0.12065 0.04679 0.00186 107.61650 24.47819 0.31388 0 0

14

23.18395 85.45948 33.67842 75.52819 36.32567 4.36836 107.61650 24.47819 0.31388

3

0 0 0 0 0 0 1.07724 24.47574 0.31388 0 0

15

23.16076 0.00299 0 0 0 0 0 0 0

4

0 0 0.74433 0.12065 0.04679 0.00186 106.53926 0.00245 0 0 0

16

0.02318 85.45648 33.67842 75.52819 36.32567 4.36836 107.61650 24.47819 0.31388

5

0.01675 61.66383 0.33027 0.05406 0.02114 0.00099 47.27126 0.00109 0 0 0

17

0.02318 85.37103 0.00002 0.00073 0.00052 0.00022 0 0 0

0.00644 23.70720 0.41408 0.06732 0.02617 0.00110 59.26799 0.00136 0 0 0

18

0 0.08546 33.67839 75.52746 36.32515 4.36814 107.61650 24.47819 0.31388 Stream Number

6

Stream Number

0 0 0 0 0 0 0 0 0 0.27340 0

19

0 0.08546 145.49042 309.35804 150.43594 14.32221 107.62726 24.47819 0.31388

7

0 0 0 0 0 0 0 0 0 0.20873 0

20

0 0.08546 144.74609 309.23739 150.38915 14.32035 0.01076 0 0

8

Table 7. Stream Flow Rates (in mol/s) for the Aromatics Complex for the HW-U-2000 Refinery, As Shown in Figure 9 9

0 0 0 0 0 0 0 0 0 0 13.62514

21

0 0 0 0 148.88526 0.10024 0 0 0

10

0 0 0 0 0 0 0 0 0 0.20873 0

22

0 0.08546 144.74609 309.23739 1.50389 14.22010 0.01076 0 0

11

14.13290 43.34339 8.37864 18.32223 8.37864 3.04385 14.83991 3.37039 0 0 0

23

0 0.08546 111.81203 233.83058 114.11079 9.95407 0.01076 0 0

12

0.36409 14.42179 6.36546 14.24242 6.95141 0.00110 31.90405 8.92746 0.31388 0 0

24

0 0.08546 0 0 0 0 0 0 0

Energy & Fuels Article

O

DOI: 10.1021/acs.energyfuels.6b00619 Energy Fuels XXXX, XXX, XXX−XXX

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biomass water investment CO2 TS&M O&M

biomass water investment CO2 TS&M O&M electricity gasoline LPG benzene toluene p-xylene o-xylene m-xylene total ($/GJ) lower bound ($/GJ) absolute difference

biomass water investment CO2 TS&M O&M electricity gasoline LPG benzene toluene p-xylene o-xylene m-xylene total ($/GJ) lower bound ($/GJ) absolute difference

8.12 0.01 17.45 0.00 4.61

U-200

U-100

8.12 0.01 22.28 0.00 5.88

5.51 0.01 17.17 0.00 4.53 2.65 −5.23 0.00 −2.71 0.00 −26.50 0.00 0.00 −4.58 −5.57 0.99

U-200

U-100

5.51 0.01 22.00 0.00 5.81 2.65 −5.23 0.00 −2.71 0.00 −26.50 0.00 0.00 1.54 0.01 1.52

9.84 0.01 17.42 0.00 4.60 2.57 −5.23 0.00 −2.71 0.00 −26.50 0.00 0.00 0.00 −1.04 1.04

U-200

9.68 0.01 22.19 0.00 5.86 2.57 −5.45 0.00 −2.37 −0.19 −26.25 0.00 0.00 6.05 4.79 1.27

U-100

8.13 0.01 12.69 0.00 3.35

U-500

5.50 0.01 12.40 0.00 3.27 2.64 −5.23 0.00 −2.71 0.00 −26.50 0.00 0.00 −10.61 −11.50 0.89

U-500

9.83 0.01 12.68 0.00 3.35 2.56 −5.23 0.00 −2.71 0.00 −26.50 0.00 0.00 −6.01 −6.83 0.82

U-500

8.13 0.01 10.06 0.00 2.66

U-1000

5.73 0.01 10.17 0.00 2.68 2.05 −5.23 0.00 −2.71 0.00 −26.50 0.00 0.00 −13.82 −14.48 0.66

U-1000

9.68 0.01 10.03 0.00 2.65 2.56 −5.45 0.00 −2.37 −0.19 −26.25 0.00 0.00 −9.34 −9.94 0.60

U-1000

8.44 0.01 8.66 0.00 2.29

U-2000

5.73 0.01 8.40 0.00 2.22 2.05 −5.23 0.00 −2.71 0.00 −26.50 0.00 0.00 −16.04 −16.62 0.58

U-2000

9.69 0.01 8.40 0.00 2.22 2.57 −5.45 0.00 −2.37 0.00 −26.48 0.00 0.00 −11.42 −12.05 0.63

U-2000

7.46 0.01 20.04 0.00 5.29

O-100

5.03 0.01 19.75 0.00 5.21 2.64 −5.32 0.00 −2.86 0.00 0.00 −22.35 0.00 2.12 0.60 1.52

O-100

9.00 0.01 20.09 0.00 5.30 2.59 −5.33 0.00 −2.88 0.00 0.00 −22.31 0.00 6.47 4.95 1.52

O-100

Table 8. Overall Cost Results (in $/GJ) for the Case Studies Are Showna

7.33 0.01 15.69 0.00 4.14

O-200

5.39 0.01 16.39 0.00 4.33 1.50 −5.24 0.00 −2.68 0.00 0.00 −22.63 0.00 −2.92 −4.08 1.16 Switchgrass

O-200

8.87 0.01 15.73 0.00 4.15 2.61 −5.24 0.00 −2.68 0.00 0.00 −22.63 0.00 0.82 −0.31 1.14 Hardwood

O-200

Corn Stover

7.91 0.01 11.88 0.00 3.14

O-500

8.07 0.03 12.49 0.00 3.30 −1.98 −5.24 0.00 −2.68 0.00 0.00 −22.63 0.00 −8.64 −9.31 0.67

O-500

9.35 0.01 11.73 0.00 3.10 1.79 −5.24 0.00 −2.68 0.00 0.00 −22.63 0.00 −4.57 −5.53 0.96

O-500

7.90 0.01 9.38 0.00 2.48

O-1000

5.39 0.01 9.11 0.00 2.41 1.50 −5.24 0.00 −2.68 0.00 0.00 −22.63 0.00 −12.13 −12.68 0.54

O-1000

9.38 0.01 9.35 0.00 2.47 1.57 −5.45 0.00 −2.36 −0.19 0.00 −22.41 0.00 −7.63 −8.42 0.79

O-1000

7.91 0.01 7.86 0.00 2.07

O-2000

5.39 0.01 7.53 0.00 1.99 1.50 −5.24 0.00 −2.68 0.00 0.00 −22.63 0.00 −14.13 −14.60 0.47

O-2000

9.45 0.01 7.82 0.00 2.06 1.48 −5.45 0.00 −2.36 −0.19 0.00 −22.41 0.00 −9.59 −10.28 0.69

O-2000

7.61 0.01 27.16 0.00 7.17

M-100

5.42 0.01 26.70 0.00 7.05 2.63 −5.45 0.00 −3.25 0.00 0.00 0.00 −22.38 10.73 8.41 2.32

M-100

9.21 0.01 27.17 0.00 7.17 2.58 −5.45 0.00 −2.36 −0.19 0.00 0.00 −23.03 15.12 13.24 1.88

M-100

7.61 0.01 21.34 0.00 5.63

M-200

8.06 0.03 22.71 0.00 6.00 −1.61 −5.45 0.00 −3.22 0.00 0.00 0.00 −22.41 4.10 2.41 1.69

M-200

9.21 0.01 21.34 0.00 5.63 2.57 −5.45 0.00 −2.36 −0.19 0.00 0.00 −23.03 7.75 6.14 1.61

M-200

8.49 0.01 16.33 0.00 4.31

M-500

5.78 0.01 16.08 0.00 4.25 0.93 −5.45 0.00 −2.36 −0.19 0.00 0.00 −23.03 −3.99 −5.07 1.08

M-500

10.07 0.02 16.21 0.00 4.28 1.16 −5.45 0.00 −2.36 −0.19 0.00 0.00 −23.03 0.69 −0.62 1.31

M-500

8.49 0.02 12.96 0.00 3.42

M-1000

5.77 0.01 12.72 0.00 3.36 0.92 −5.45 0.00 −2.36 −0.19 0.00 0.00 −23.03 −8.26 −9.02 0.76

M-1000

10.19 0.02 12.90 0.00 3.41 0.99 −5.45 0.00 −2.36 −0.19 0.00 0.00 −23.03 −3.54 −4.60 1.06

M-1000

8.49 0.02 10.72 0.00 2.83

M-2000

5.87 0.01 10.49 0.00 2.77 0.91 −5.24 0.00 −2.69 0.00 0.00 0.00 −23.25 −11.12 −11.72 0.60

M-2000

10.23 0.02 10.69 0.00 2.82 0.91 −5.45 0.00 −2.36 −0.19 0.00 0.00 −23.03 −6.36 −7.31 0.95

M-2000

Energy & Fuels Article

DOI: 10.1021/acs.energyfuels.6b00619 Energy Fuels XXXX, XXX, XXX−XXX

Article

shown, the more profitable the refinery is. The value shown in Table 8 takes the (i) feedstock costs, (ii) CO2 sequestration costs, (iii) investment costs, (iv) operating and maintenance (O&M) costs, (v) electricity costs or revenue, (vi) revenue from selling LPG, (vii) revenue from liquid products, and (viii) revenue from aromatics into account. The values shown in Table 8 are normalized with respect to the total energy (in GJ) of products produced. The absolute difference between the lower bound and the upper bound of the solutions obtained is shown; this difference represents how close the solutions obtained are to the mathematically best possible solution. Across all three feedstock types, the unrestricted refineries are the most profitable. They are followed by the o-xylene refineries and the m-xylene refineries, respectively. The unrestricted refineries mainly produce p-xylene, with smaller amounts of benzene and gasoline also produced. Gasoline byproduct is inevitable in each of the case studies since the methanol-to-aromatics reaction produces some C6+ paraffinic components. The Cyclar process is utilized to convert all low-quality LPG into high-value aromatics; therefore, no LPG is output from the refineries. Electricity is input into every BTA refinery except for the HW-O-500 and HW-M-200 refineries, for which output electricity is sold to the grid. The m-xylene refineries are the least profitable refineries because of the increased investment cost with the UOP MX Sorbex process. The hardwood aromatics refineries are the most profitable refineries across all case studies. The switchgrass refineries are the second most profitable refineries, with the corn stover refineries being the least profitable. This trend follows the costs of the individual feedstocks: with hardwood ($70/dry metric ton) being the least expensive, followed by switchgrass ($100/dry metric ton), and finally corn stover ($120/dry metric ton). The refineries at small scales (≤500 MT/day) represent pilot plant scale refineries. All of the 100 t per day case studies are unprofitable. This is in addition to the CS-U-200, CS-O-200, CS-M-200, CS-M-500, HW-M-200, and SG-M-200 refineries, which are also unprofitable. As Table 8 illustrates, the investment cost represents the largest expenditure associated with each of the hardwood aromatics refineries. As Table 8 shows and is made more explicit in Table 9, the o-xylene refineries always have the lowest required investment cost across similar capacity BTA plants. This is also clear from comparing the values across similar capacity plants in the Investment row in Table 8. However, despite the increased investment costs associated with the unrestricted refineries, they are still more profitable than the o-xylene refineries because of the higher value of p-xylene compared to o-xylene. As an example, Table 8 shows that the investment cost for the HW-O-2000 refinery ($7.53/GJ) is $0.87/GJ less expensive than the HW-U-2000 refinery ($8.40/GJ). However, the revenue from selling p-xylene in the HW-U-2000 refinery ($26.50/GJ) is $3.87/GJ higher than from selling o-xylene in the HW-O-2000 refinery ($22.63/GJ). This relative difference in the selling prices of the two C8 isomers is the main driver in determining the relative profitability in the unrestricted and o-xylene refineries. The HW-U-2000 case study has an overall profit of $16.04/GJ, the SG-U-2000 refinery has an overall profit of $13.11/GJ, and the CS-U-2000 case study has an overall profit of $11.42/GJ. 3.3. Investment Costs. The total capital investment cost, or the fixed capital investment (FCI), is shown in Table 9 for each case study. The breakdown for each section of the BTA refinery is also shown; it consists of the syngas generation, syngas cleaning, hydrocarbon production, hydrocarbon upgrading, chemicals production, hydrogen and oxygen production, heat and power integration, and wastewater treatment sections. In the unrestricted corn stover case studies, the largest contributor to the investment cost is the syngas generation section. This section constitutes between 23 and 29% of the total investment cost across these case studies. Likewise, the syngas generation section is also the largest contributor to the investment cost in all of the unrestricted switchgrass refineries. This section constitutes between 24 and 30% of the total investment cost across the case studies. The largest contributor to the total investment cost for the unrestricted hardwood case studies is the chemicals production section at scales less than or equal to 500 t per day.

a Contributions to the total costs (in $/GJ) of the BTA refineries come from biomass, water, CO2 transportation/storage/monitoring, investment, input electricity, and operating and maintenance. The more negative this value is the more profitable a refinery becomes. The overall profit can be determined by taking the negative of the “total ($/GJ)” value. The lower bound values are reported along with the absolute difference between the upper bound and the lower bound.

0.83 −5.45 0.00 −2.36 −0.19 0.00 0.00 −23.03 −8.14 −8.89 0.75 0.83 −5.45 0.00 −2.36 −0.19 0.00 0.00 −23.03 −5.31 −6.25 0.93

M-1000 M-500

0.83 −5.45 0.00 −2.36 −0.19 0.00 0.00 −23.03 −1.06 −2.26 1.20 2.52 −5.45 0.00 −2.36 −0.19 0.00 0.00 −23.03 6.08 4.37 1.71

M-200 M-100

2.52 −5.45 0.00 −2.36 −0.19 0.00 0.00 −23.03 13.45 11.15 2.30 1.44 −5.24 0.00 −2.68 0.00 0.00 −22.63 0.00 −11.25 −11.85 0.60

O-2000 O-1000

1.44 −5.24 0.00 −2.68 0.00 0.00 −22.63 0.00 −9.34 −10.03 0.69 1.44 −5.24 0.00 −2.68 0.00 0.00 −22.63 0.00 −6.16 −6.98 0.82

O-500 O-200

2.55 −5.25 0.00 −2.68 0.00 0.00 −22.62 0.00 −0.82 −2.08 1.26 2.53 −5.35 0.00 −2.94 0.00 0.00 −22.22 0.00 4.81 3.31 1.50 1.93 −5.23 0.00 −2.71 0.00 −26.50 0.00 0.00 −13.11 −13.72 0.61 2.51 −5.23 0.00 −2.71 0.00 −26.50 0.00 0.00 −11.07 −11.76 0.69 2.51 −5.23 0.00 −2.71 0.00 −26.50 0.00 0.00 −7.75 −8.60 0.84 2.51 −5.23 0.00 −2.71 0.00 −26.50 0.00 0.00 −1.74 −2.79 1.05 2.52 −5.23 0.00 −2.71 0.00 −26.50 0.00 0.00 4.37 2.91 1.46

Switchgrass

O-100 U-2000 U-1000 U-500 U-200 U-100

Table 8. continued

electricity gasoline LPG benzene toluene p-xylene o-xylene m-xylene total ($/GJ) lower bound ($/GJ) absolute difference

M-2000

Energy & Fuels

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DOI: 10.1021/acs.energyfuels.6b00619 Energy Fuels XXXX, XXX, XXX−XXX

R

a

76.63 52.56 25.27 38.54 70.63 32.51 2.99 9.48 308.61

U-200

U-100

47.32 33.70 16.52 24.56 44.33 22.68 1.92 5.99 197.02

68.21 49.50 25.09 38.56 70.63 39.23 3.04 9.45 303.69

U-200

U-100

41.80 31.96 16.31 24.57 44.33 27.62 1.94 6.04 194.58

74.96 52.94 26.16 38.55 70.63 32.37 3.13 9.42 308.15

U-200

45.74 33.82 16.99 24.54 44.07 23.06 2.01 6.00 196.24

All values are shown in MM$.

syngas generation syngas cleanup HC production HC upgrading chemicals production H2/O2 production H&P integration wastewater treatment total (MM$)

syngas generation syngas cleanup HC production HC upgrading chemicals production H2/O2 production H&P integration wastewater treatment total (MM$)

syngas generation syngas cleanup HC production HC upgrading chemicals production H2/O2 production H&P integration wastewater treatment total (MM$)

U-100

146.28 94.89 44.82 69.90 130.81 51.86 5.53 16.96 561.05

U-500

129.65 88.30 44.48 69.93 130.81 62.62 5.64 17.10 548.54

U-500

143.12 95.57 46.38 69.92 130.81 51.72 5.79 17.23 560.54

U-500

244.54 148.62 69.21 109.68 208.58 73.95 8.80 26.59 889.97

U-1000

220.50 137.47 68.82 109.75 208.58 87.25 39.29 27.38 899.04

U-1000

237.75 149.30 71.63 109.59 207.38 74.77 9.38 26.91 886.71

U-1000

455.07 234.94 121.34 193.52 332.70 101.42 50.18 42.96 1532.13

U-2000

397.75 214.62 120.56 193.67 332.70 125.14 57.91 43.00 1485.35

U-2000

432.51 234.98 127.64 193.08 334.25 106.70 14.84 42.21 1486.20

U-2000

45.83 33.44 16.87 24.55 23.96 24.68 2.18 5.70 177.21

O-100

40.87 31.72 16.58 24.56 24.10 28.89 2.22 5.74 174.67

O-100

45.00 33.71 17.63 24.56 24.06 24.71 2.28 5.74 177.67

O-100

74.36 52.23 26.52 38.53 37.82 35.69 3.57 8.91 277.62

O-200

67.81 49.39 25.04 38.55 37.83 39.65 22.36 9.36 289.99 Switchgrass

O-200

73.17 52.66 27.67 38.53 37.83 35.63 3.71 8.97 278.18 Hardwood

O-200

Corn Stover

144.90 94.61 44.88 69.90 67.87 53.25 33.27 16.79 525.46

O-500

131.23 93.49 48.56 70.05 67.87 45.58 71.41 24.35 552.54

O-500

141.05 95.12 46.86 69.91 67.87 54.29 26.87 16.75 518.71

O-500

241.06 148.16 69.28 109.67 105.89 76.00 52.73 26.57 829.35

O-1000

213.60 136.57 68.51 109.73 105.89 90.40 54.46 26.56 805.71

O-1000

234.39 148.92 71.87 109.61 105.72 77.02 53.08 26.26 826.87

O-1000

440.83 232.68 121.33 193.51 165.54 108.49 86.04 41.32 1389.73

O-2000

388.53 212.54 119.46 193.63 165.54 129.18 80.82 41.71 1331.42

O-2000

428.01 234.34 128.43 193.27 165.27 109.30 82.85 41.83 1383.29

O-2000

46.05 33.44 16.61 24.53 87.40 24.21 2.09 5.90 240.24

M-100

41.72 31.89 16.28 24.57 85.90 27.80 1.97 5.98 236.12

M-100

45.18 33.69 17.27 24.54 87.40 24.14 2.17 5.94 240.33

M-100

75.00 52.20 25.60 38.49 139.12 34.46 3.32 9.24 377.43

M-200

68.12 52.35 27.48 38.60 137.16 28.55 36.63 12.82 401.70

M-200

73.61 52.60 26.59 38.50 139.12 34.37 3.45 9.29 377.53

M-200

148.47 95.35 44.80 69.84 257.22 49.37 39.51 17.64 722.19

M-500

130.92 88.81 44.68 69.88 257.22 60.99 41.05 17.68 711.23

M-500

144.06 95.77 46.27 69.86 257.22 50.31 35.93 17.24 716.64

M-500

250.02 149.35 69.18 109.59 409.51 70.36 61.11 27.37 1146.49

M-1000

221.38 137.51 68.79 109.65 409.51 86.63 63.54 27.70 1124.72

M-1000

243.58 150.23 71.44 109.61 409.51 70.81 58.20 27.32 1140.70

M-1000

456.12 234.93 121.20 193.26 652.07 100.49 95.83 42.96 1896.86

M-2000

400.74 215.57 121.23 193.68 657.66 123.57 99.70 43.66 1855.81

M-2000

445.70 236.83 126.70 193.29 652.07 101.42 92.12 43.43 1891.56

M-2000

Table 9. Investment Costs of the Syngas Generation, Syngas Cleaning, Hydrocarbon Production, Hydrocarbon Upgrading, Chemicals Production, Hydrogen/Oxygen Production, Heat and Power Integration, and Wastewater Treatment Sections of the BTA Refinery Are Showna

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Article

Energy & Fuels At scales larger than 500 t per day, the syngas generation section is the largest contributor to the investment for the unrestricted hardwood case studies. In all of the o-xylene case studies, the largest contributor to the investment cost is the syngas generation section. This is because the chemicals production section does not include any expensive adsorptive technologies. The chemicals production section is the largest contributor in the m-xylene case studies because of the high capital costs associated with m-xylene separation. The investment costs across feedstock types are relatively similar for each individual refinery scale and product type. The value of the investment costs ranges from $174 to $240 MM for the 100 MT/day plants, $277 to $402 MM for the 200 MT/day plants, $518 to $722 MM for the 500 MT/day plants, $806 to $1147 MM for the 1000 MT/day plants, and $1331 to $1896 MM for the 2000 MT/day plants. The o-xylene case studies are consistently the least expensive refineries because of the absence of adsorptive technologies. The m-xylene case studies are consistently the most expensive refineries because of the large cost associated with the MX Sorbex unit. 3.4. Discounted Cash Flow Analysis. A discounted cash flow analysis is carried out for all case studies investigated. Our previous studies have described the main parameters and methodology used in this analysis in detail.95,108 We use the same methodology in this paper to calculate the net present value (NPV) for each biomass to aromatics refinery. The main parameters, as well as a description of our methodology,

is included in Appendix A.2 (Supporting Information). For more information on net present value calculations, the reader is directed to the following textbook.171 The net present values for the corn stover, hardwood, and switchgrass refineries are shown in Figures 10, 11, and 12, respectively. Since the BTA refineries at small scales (≤500 MT/day) represent pilot plants, they just barely break-even or do not return the initial investment. The hardwood case studies have the highest net present values, followed by the switchgrass case studies and finally the corn stover refineries. The HW-U-2000 refinery has a net present value of over $1200 MM, the SG-U-2000 refinery has a NPV of over $750 MM, and the CS-U-2000 refinery has a NPV of over $550 MM dollars. Table 10 illustrates the amount of time it takes to pay back the initial investment of the plants. The most profitable biomass to aromatics refineries take roughly 9 years to pay back at the largest scales. 3.5. Material and Energy Balances. Table 11 shows the material balances for the BTA refineries, while the energy balances are shown in Table 12. The amount of biomass input into the refinery is shown in dry tons per hour, while the freshwater input is shown in thousand barrels per day. The amount of LPG and gasoline output are also shown in thousand barrels per day, the amount of aromatics output is shown is metric tons per day, and the amount of vented and sequestered CO2 is shown in metric tons per hour.

Figure 10. Net present values for the corn stover case studies.

Figure 11. Net present values for the hardwood case studies. S

DOI: 10.1021/acs.energyfuels.6b00619 Energy Fuels XXXX, XXX, XXX−XXX

Article

Energy & Fuels

Figure 12. Net present values for the switchgrass case studies.

Table 10. Payback Period in Years for Each Case Study U-100 corn stover hardwood switchgrass

U-200

U-500

U-1000

U-2000

29

26 13 18

13 9 11

O-100

O-200

O-500

O-1000

O-2000

13 21

15 9 12

M-100

M-200

M-500

M-1000

M-2000 17

to photosynthesis or soil storage. The amount of atmospheric CO2 that is captured during photosynthesis in part (e) is based on the carbon content of the biomass.9 Perennial crops are assumed to store 0.3 g of carbon in degraded lands per gram of as-received biomass over the lifetime of the crop.9 Using information on the emissions avoided from chemicals (GHGAC), liquid fuels (GHGAF), and electricity (GHGAE) that was described earlier, the relative life cycle emissions to typical fossil fuel based process are compared. The greenhouse gas index (GHGI) is calculated by dividing the life cycle greenhouse gas emissions (LGHG) by the summation of the GHGAC, GHGAF, and GHGAE. A GHGI of less than 1 has superior life cycle emissions compared to fossil fuel based processes, while a GHGI that is negative indicates processes that absorb more CO2 than is released by the refinery. As Table 14 indicates, all biomass to aromatics refineries achieve significant reduction in GHG emissions compared to fossil fuel based processes. For corn stover and hardwood refineries, this reduction in life cycle GHG emissions is between 89 and 90% and 80−85%, respectively. However, the most environmentally beneficial biomass aromatics use perennial crops, which have GHGI’s between −0.8 and −0.73. These are net-negative processes due to the additional CO2 captured and stored in the soil. 3.7. Parametric Analysis. The impact of higher or lower feedstock costs on the overall profitability of the unrestricted 2000 MT/day aromatics refineries is examined. The net present values illustrated in Figures 10, 11, and 12 are based on corn stover, hardwood, and switchgrass prices of $120, $70, and $100 per dry metric ton, respectively. The net present value is recalculated for corn stover prices of $100 and $140 per dry metric ton, hardwood prices of $60 and $80 per dry metric ton, and switchgrass prices of $80 and $120 per dry metric ton. The values are then compared with the nominal results presented in Figures 10, 11, and 12. This comparison is explicitly shown in Table 15. For a hardwood price of $60 the NPV of the U-2000 aromatics refinery is $1340 MM. The NPV of the HW-U-2000 refinery decreases to $1102 MM when the hardwood price increases to $80. Similarly, the NPVs of the corn stover U-2000 refineries decrease from $786 MM to $310 MM when the price increases from $100 to $140 per dry metric ton. The switchgrass refineries have NPVs between $532 and $1023 MM when the switchgrass price ranges from $80 to $120.

Across all feedstock types, the unrestricted 2000 MT/day refineries produce roughly 450 thousand metric tons of p-xylene per year. These case studies also produce benzene and byproduct gasoline. The CS-U-2000 case study requires 278 dt/h of corn stover, the HW-U-2000 case study requires 282 dt/h of hardwood, and the SG-U-2000 case study requires roughly 291 dt/h of switchgrass. None of the case studies sequestered CO2 produced within the refinery. The reader is referred to Table 11 for the amounts of biomass required for the o-xylene and m-xylene case studies as well as the corresponding product outputs. The contributions to the overall energy balance for all case studies is illustrated in Table 12. The energy inputs include electricity and biomass, while the energy outputs include the aromatics, LPG, gasoline, and electricity. If electricity is input into (output from) the refinery, it is shown as a positive (negative) value. The overall energy efficiency is also calculated. The unrestricted corn stover refineries have energy efficiencies between 66 and 67%, the unrestricted hardwood refineries have energy efficiencies between 63 and 65%, and the unrestricted switchgrass refineries have energy efficiencies between 63 and 65%. 3.6. Carbon and Greenhouse Gas Balances. Table 13 shows the carbon balances for each of the case studies. The only carbon input into the BTA refineries is biomass; the carbon outputs include the aromatics, gasoline, LPG, vented CO2, and sequestered CO2. The amount of carbon input into the refineries by air is considered negligible and is therefore not shown in Table 13. Carbon conversion is calculated for all refineries by summing the carbon output as products and dividing it by the carbon input into the refinery. The unrestricted corn stover refineries have carbon conversions of approximately 55−56%, the unrestricted hardwood refineries have carbon conversions between 51 and 54%, and the unrestricted switchgrass refineries have carbon conversions between 53 and 56%. Table 14 shows the total greenhouse gas emissions (in kg CO2 equivalent/s) for the inputs and outputs of the BTA refineries. As mentioned previously, an environmental constraint imposed in the process synthesis model forces the BTA refineries to have at least 50% reduction in life cycle greenhouse gas emissions from current fossil-fuel based processes. The life cycle GHG emissions (GHG) are calculated by summing the respective emissions rates from each stage of production, which include (a) the acquisition and transportation of the biomass feed, (b) the transportation and use of the aromatics, gasoline, and LPG (c) venting of CO2, (d) the transportation and sequestration of CO2, and (e) atmospheric sequestration of CO2 during growth of biomass due T

DOI: 10.1021/acs.energyfuels.6b00619 Energy Fuels XXXX, XXX, XXX−XXX

U

a

27.96 0.00 0.36 15.63 0.00 136.66 0.00 0.00 1.38 0.00 21.24

U-200

U-100

13.98 0.00 0.18 7.82 0.00 68.33 0.00 0.00 0.70 0.00 10.61

27.07 0.00 0.36 15.63 0.00 136.66 0.00 0.00 0.83 0.00 23.16

U-200

U-100

13.55 0.00 0.18 7.82 0.00 68.33 0.00 0.00 0.41 0.00 11.60

28.22 0.00 0.36 15.63 0.00 136.66 0.00 0.00 1.54 0.00 21.57

U-200

13.89 0.00 0.19 6.83 0.62 67.66 0.00 0.00 0.76 0.00 10.42

U-100

69.95 0.00 0.90 39.08 0.00 341.65 0.00 0.00 3.49 0.00 53.18

U-500

67.65 0.00 0.90 39.08 0.00 341.65 0.00 0.00 2.06 0.00 57.87

U-500

70.50 0.00 0.90 39.08 0.00 341.65 0.00 0.00 3.78 0.00 53.83

U-500

139.88 0.00 1.80 78.17 0.00 683.30 0.00 0.00 6.98 0.00 106.34

U-1000

140.75 0.00 1.80 78.17 0.00 683.30 0.00 0.00 4.65 0.00 125.83

U-1000

138.75 0.00 1.88 68.32 6.19 676.62 0.00 0.00 7.42 0.00 103.99

U-1000

290.63 0.00 3.60 156.33 0.00 1366.59 0.00 0.00 15.07 0.00 231.40

U-2000

281.96 0.00 3.60 156.33 0.00 1366.59 0.00 0.00 9.35 0.00 252.50

U-2000

277.79 0.00 3.75 136.94 0.00 1365.22 0.00 0.00 14.87 0.00 208.50

U-2000

12.84 0.00 0.18 8.47 0.00 0.00 67.14 0.00 0.56 0.00 8.66

O-100

12.35 0.00 0.18 8.25 0.00 0.00 67.51 0.00 0.29 0.00 9.40

O-100

12.90 0.00 0.18 8.32 0.00 0.00 67.40 0.00 0.62 0.00 8.71

O-100

25.23 0.00 0.36 15.48 0.00 0.00 136.68 0.00 1.09 0.00 16.54

O-200

26.51 0.00 0.36 15.46 0.00 0.00 136.72 0.00 0.78 0.00 22.14 Switchgrass

O-200

25.44 0.00 0.36 15.46 0.00 0.00 136.72 0.00 1.22 0.00 16.79 Hardwood

O-200

Corn Stover

68.08 0.00 0.90 38.64 0.00 0.00 341.80 0.00 3.31 0.00 49.97

O-500

99.16 0.00 0.90 38.64 0.00 0.00 341.80 0.00 7.15 0.00 116.16

O-500

67.01 0.00 0.90 38.64 0.00 0.00 341.80 0.00 3.35 0.00 47.83

O-500

136.01 0.00 1.80 77.29 0.00 0.00 683.60 0.00 6.51 0.00 99.70

O-1000

132.53 0.00 1.80 77.29 0.00 0.00 683.60 0.00 3.87 0.00 110.64

O-1000

134.53 0.00 1.88 68.05 5.94 0.00 677.12 0.00 7.04 0.00 96.75

O-1000

272.38 0.00 3.61 154.58 0.00 0.00 1367.20 0.00 13.30 0.00 200.01

O-2000

265.16 0.00 3.61 154.58 0.00 0.00 1367.20 0.00 7.78 0.00 221.46

O-2000

270.98 0.00 3.75 136.11 11.88 0.00 1354.24 0.00 14.21 0.00 196.79

O-2000

13.10 0.00 0.19 6.82 0.60 0.00 0.00 67.69 0.64 0.00 9.13

M-100

13.33 0.00 0.19 9.37 0.00 0.00 0.00 65.78 0.39 0.00 11.21

M-100

13.21 0.00 0.19 6.82 0.60 0.00 0.00 67.69 0.70 0.00 9.26

M-100

26.20 0.00 0.38 13.64 1.21 0.00 0.00 135.38 1.27 0.00 18.25

M-200

39.65 0.00 0.38 18.57 0.00 0.00 0.00 131.72 2.61 0.00 46.45

M-200

26.43 0.00 0.38 13.64 1.21 0.00 0.00 135.38 1.41 0.00 18.53

M-200

73.04 0.00 0.94 34.09 3.01 0.00 0.00 338.46 3.88 0.00 58.61

M-500

71.08 0.00 0.94 34.09 3.01 0.00 0.00 338.46 2.47 0.00 64.33

M-500

72.21 0.00 0.94 34.09 3.01 0.00 0.00 338.46 4.03 0.00 56.87

M-500

146.10 0.00 1.88 68.17 6.02 0.00 0.00 676.92 7.84 0.00 117.26

M-1000

141.93 0.00 1.88 68.17 6.02 0.00 0.00 676.92 4.94 0.00 128.22

M-1000

146.09 0.00 1.88 68.17 6.02 0.00 0.00 676.92 8.35 0.00 116.60

M-1000

The inputs to the BTA refinery are biomass and water, while the outputs include gasoline, LPG, benzene, toluene, p-xylene, o-xylene, m-xylene, sequestered CO2, and vented CO2.

biomass (dt/h) LPG (kBD) gasoline (kBD) benzene [MT/day] toluene [MT/day] p-xylene [MT/day] o-xylene [MT/day] m-xylene [MT/day] water (kBD) seq CO2 (MT/h) vented CO2 (MT/h)

biomass (dt/h) LPG (kBD) gasoline (kBD) benzene [MT/day] toluene [MT/day] p-xylene [MT/day] o-xylene [MT/day] m-xylene [MT/day] water (kBD) seq CO2 (MT/h) vented CO2 (MT/h)

biomass (dt/h) LPG (kBD) gasoline (kBD) benzene [MT/day] toluene [MT/day] p-xylene [MT/day] o-xylene [MT/day] m-xylene [MT/day] water (kBD) seq CO2 (MT/h) vented CO2 (MT/h)

Table 11. Overall Material Balances for the Case Studies Are Showna

292.32 0.00 3.75 136.35 12.05 0.00 0.00 1353.84 15.74 0.00 234.72

M-2000

288.85 0.00 3.60 155.25 0.00 0.00 0.00 1366.93 10.12 0.00 265.27

M-2000

293.29 0.00 3.75 136.35 12.05 0.00 0.00 1353.84 16.72 0.00 235.10

M-2000

Energy & Fuels Article

DOI: 10.1021/acs.energyfuels.6b00619 Energy Fuels XXXX, XXX, XXX−XXX

V

U-200

134.84 22.92 0.00 7.36 0.00 65.32 0.00 0.00 12.36 64.95

U-100

67.40 11.46 0.00 3.68 0.00 32.66 0.00 0.00 6.20 64.95

134.15 22.92 0.00 7.36 0.00 65.32 0.00 0.00 13.01 64.96

67.13 11.46 0.00 3.68 0.00 32.66 0.00 0.00 6.53 64.89

U-200

U-100

U-200

132.09 22.92 0.00 7.36 0.00 65.32 0.00 0.00 12.62 66.07

65.00 11.95 0.00 3.21 0.29 32.34 0.00 0.00 6.31 67.04

U-100

U-500

337.30 57.31 0.00 18.39 0.00 163.31 0.00 0.00 30.87 64.92

U-500

335.27 57.31 0.00 18.39 0.00 163.31 0.00 0.00 32.49 64.99

U-500

329.96 57.31 0.00 18.39 0.00 163.31 0.00 0.00 31.52 66.12

674.53 114.61 0.00 36.78 0.00 326.62 0.00 0.00 61.71 64.93

U-1000

697.56 114.61 0.00 36.78 0.00 326.62 0.00 0.00 50.28 63.92

U-1000

649.39 119.50 0.00 32.14 2.94 323.43 0.00 0.00 62.96 67.10

U-1000

U-2000

1401.48 229.22 0.00 73.55 0.00 653.24 0.00 0.00 94.93 63.89

U-2000

1397.40 229.22 0.00 73.55 0.00 653.24 0.00 0.00 100.79 63.81

U-2000

1300.13 239.00 0.00 64.43 0.00 652.59 0.00 0.00 126.12 67.03

61.90 11.72 0.00 3.99 0.00 0.00 32.09 0.00 6.21 70.18

O-100

61.22 11.65 0.00 3.88 0.00 0.00 32.27 0.00 6.50 70.58

O-100

60.37 11.67 0.00 3.91 0.00 0.00 32.22 0.00 6.36 71.63

O-100

O-500

O-500

121.64 22.98 0.00 7.28 0.00 0.00 65.33 0.00 12.53 71.25

O-200 328.27 57.44 0.00 18.18 0.00 0.00 163.38 0.00 17.71 69.08

O-500

131.41 491.45 22.98 57.44 0.00 0.00 7.27 18.18 0.00 0.00 0.00 0.00 65.35 163.38 0.00 0.00 7.38 −24.35 68.88 53.59 Switchgrass

O-200

119.05 313.62 22.98 57.44 0.00 0.00 7.27 18.18 0.00 0.00 0.00 0.00 65.35 163.38 0.00 0.00 12.83 22.02 72.49 71.21 Hardwood

O-200

Corn Stover

655.89 114.88 0.00 36.36 0.00 0.00 326.77 0.00 35.49 69.14

O-1000

656.84 114.88 0.00 36.36 0.00 0.00 326.77 0.00 36.91 68.90

O-1000

629.66 119.50 0.00 32.02 2.82 0.00 323.67 0.00 38.62 71.53

O-1000

O-2000

1313.48 229.76 0.00 72.73 0.00 0.00 653.53 0.00 70.89 69.06

O-2000

1314.15 229.76 0.00 72.73 0.00 0.00 653.53 0.00 73.85 68.88

O-2000

1268.25 239.00 0.00 64.04 5.64 0.00 647.34 0.00 72.60 71.30

63.18 11.95 0.00 3.21 0.29 0.00 0.00 32.36 6.20 68.90

M-100

66.07 11.95 0.00 4.41 0.00 0.00 0.00 31.44 6.47 65.89

M-100

61.84 11.95 0.00 3.21 0.29 0.00 0.00 32.36 6.33 70.11

M-100

M-200

126.35 23.90 0.00 6.42 0.57 0.00 0.00 64.71 12.39 68.91

M-200

196.50 23.90 0.00 8.74 0.00 0.00 0.00 62.96 −7.93 52.69

M-200

123.68 23.90 0.00 6.42 0.57 0.00 0.00 64.71 12.66 70.12

352.21 59.75 0.00 16.04 1.43 0.00 0.00 161.79 10.19 65.95

M-500

352.29 59.75 0.00 16.04 1.43 0.00 0.00 161.79 11.38 65.72

M-500

337.98 59.75 0.00 16.04 1.43 0.00 0.00 161.79 14.23 67.86

M-500

704.52 119.50 0.00 32.08 2.86 0.00 0.00 323.57 20.41 65.94

M-1000

703.42 119.50 0.00 32.08 2.86 0.00 0.00 323.57 22.63 65.84

M-1000

683.75 119.50 0.00 32.08 2.86 0.00 0.00 323.57 24.33 67.51

M-1000

1409.63 239.00 0.00 64.15 5.71 0.00 0.00 647.15 40.81 65.91

M-2000

1431.57 229.57 0.00 73.04 0.00 0.00 0.00 653.40 44.98 64.75

M-2000

1372.68 239.00 0.00 64.15 5.71 0.00 0.00 647.15 44.96 67.44

M-2000

The energy inputs come from biomass or electricity, while the energy outputs are the aromatics, gasoline, LPG, or electricity. Electricity is denoted as a positive value if it is input into the system and as a negative value if it is output from the system. The overall efficiency (in percent) of the refineries is calculated by dividing the energy outputs by the energy inputs.

a

biomass gasoline LPG benzene toluene p-xylene o-xylene m-xylene electricity efficiency (%)

biomass gasoline LPG benzene toluene p-xylene o-xylene m-xylene electricity efficiency (%)

biomass gasoline LPG benzene toluene p-xylene o-xylene m-xylene electricity efficiency (%)

Table 12. Overall Energy Balances (in MW) for the Case Studies Are Showna

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Article

Energy & Fuels Table 13. Carbon Accounting (in kg/s) Is Shown for the Case Studiesa Corn Stover biomass gasoline LPG benzene toluene p-xylene o-xylene m-xylene vented CO2 seq CO2 conversion (%)

U-100

U-200

U-500

U-1000

U-2000

O-100

O-200

O-500

O-1000

O-2000

M-100

M-200

M-500

M-1000

M-2000

1.80 0.23 0.00 0.07 0.01 0.71 0.00 0.00 0.79 0.00 56.25

3.67 0.43 0.00 0.17 0.00 1.43 0.00 0.00 1.64 0.00 55.44

9.16 1.09 0.00 0.42 0.00 3.58 0.00 0.00 4.08 0.00 55.49

18.02 2.27 0.00 0.73 0.07 7.09 0.00 0.00 7.88 0.00 56.30

36.08 4.53 0.00 1.46 0.00 14.30 0.00 0.00 15.81 0.00 56.24

1.68 0.22 0.00 0.09 0.00 0.00 0.71 0.00 0.66 0.00 60.63

3.30 0.44 0.00 0.17 0.00 0.00 1.43 0.00 1.27 0.00 61.51

8.70 1.09 0.00 0.41 0.00 0.00 3.58 0.00 3.63 0.00 58.37

17.48 2.27 0.00 0.73 0.06 0.00 7.09 0.00 7.33 0.00 58.07

35.20 4.53 0.00 1.45 0.13 0.00 14.19 0.00 14.92 0.00 57.66

1.72 0.23 0.00 0.07 0.01 0.00 0.00 0.71 0.70 0.00 59.12

3.43 0.45 0.00 0.15 0.01 0.00 0.00 1.42 1.40 0.00 59.12

9.38 1.13 0.00 0.36 0.03 0.00 0.00 3.55 4.31 0.00 54.09

18.98 2.27 0.00 0.73 0.06 0.00 0.00 7.09 8.84 0.00 53.47

38.10 4.53 0.00 1.46 0.13 0.00 0.00 14.18 17.82 0.00 53.27

Hardwood biomass gasoline LPG benzene toluene p-xylene o-xylene m-xylene vented CO2 seq CO2 conversion (%)

U-100

U-200

U-500

U-1000

U-2000

O-100

O-200

O-500

O-1000

O-2000

M-100

M-200

M-500

M-1000

M-2000

1.89 0.22 0.00 0.08 0.00 0.72 0.00 0.00 0.88 0.00 53.81

3.77 0.43 0.00 0.17 0.00 1.43 0.00 0.00 1.76 0.00 53.86

9.43 1.09 0.00 0.42 0.00 3.58 0.00 0.00 4.39 0.00 53.88

19.62 2.17 0.00 0.83 0.00 7.16 0.00 0.00 9.54 0.00 51.79

39.31 4.34 0.00 1.67 0.00 14.32 0.00 0.00 19.14 0.00 51.71

1.72 0.22 0.00 0.09 0.00 0.00 0.71 0.00 0.71 0.00 58.99

3.70 0.44 0.00 0.17 0.00 0.00 1.43 0.00 1.68 0.00 54.98

13.83 1.09 0.00 0.41 0.00 0.00 3.58 0.00 8.81 0.00 36.75

18.48 2.18 0.00 0.83 0.00 0.00 7.16 0.00 8.39 0.00 55.00

36.97 4.35 0.00 1.65 0.00 0.00 14.32 0.00 16.79 0.00 54.98

1.86 0.23 0.00 0.10 0.00 0.00 0.00 0.69 0.85 0.00 54.65

5.53 0.45 0.00 0.20 0.00 0.00 0.00 1.38 3.52 0.00 36.75

9.91 1.13 0.00 0.36 0.03 0.00 0.00 3.55 4.88 0.00 51.20

19.79 2.27 0.00 0.73 0.06 0.00 0.00 7.09 9.72 0.00 51.28

40.27 4.35 0.00 1.66 0.00 0.00 0.00 14.32 20.11 0.00 50.47

U-100

U-200

U-500

U-1000

U-2000

O-100

O-200

O-500

O-1000

O-2000

M-100

M-200

M-500

M-1000

M-2000

1.82 0.22 0.00 0.08 0.00 0.72 0.00 0.00 0.80 0.00 55.84

3.64 0.43 0.00 0.17 0.00 1.43 0.00 0.00 1.61 0.00 55.82

9.11 1.09 0.00 0.42 0.00 3.58 0.00 0.00 4.03 0.00 55.78

18.22 2.17 0.00 0.83 0.00 7.16 0.00 0.00 8.06 0.00 55.79

37.85 4.34 0.00 1.67 0.00 14.32 0.00 0.00 17.54 0.00 53.70

1.67 0.22 0.00 0.09 0.00 0.00 0.70 0.00 0.66 0.00 60.77

3.29 0.44 0.00 0.17 0.00 0.00 1.43 0.00 1.25 0.00 61.87

8.87 1.09 0.00 0.41 0.00 0.00 3.58 0.00 3.79 0.00 57.31

17.71 2.18 0.00 0.83 0.00 0.00 7.16 0.00 7.56 0.00 57.37

35.47 4.35 0.00 1.65 0.00 0.00 14.32 0.00 15.16 0.00 57.30

1.71 0.23 0.00 0.07 0.01 0.00 0.00 0.71 0.69 0.00 59.47

3.41 0.45 0.00 0.15 0.01 0.00 0.00 1.42 1.38 0.00 59.48

9.51 1.13 0.00 0.36 0.03 0.00 0.00 3.55 4.44 0.00 53.34

19.03 2.27 0.00 0.73 0.06 0.00 0.00 7.09 8.89 0.00 53.33

38.07 4.53 0.00 1.46 0.13 0.00 0.00 14.18 17.79 0.00 53.31

Switchgrass biomass gasoline LPG benzene toluene p-xylene o-xylene m-xylene vented CO2 seq CO2 conversion (%)

a Biomass is the carbon input, while the carbon outputs are the aromatics, the liquid products, the LPG byproduct, vented CO2, or sequestered (Seq.) CO2. Carbon conversion for each case study is calculated by dividing the total carbon exiting as either aromatics, liquid product, LPG by the total carbon input into the refinery.

4. CONCLUSIONS

The most profitable BTA refineries had net present values greater than $1200 MM dollars with payback periods less than 10 years. The required investment cost for a plant producing approximately half a million metric tons of aromatics per year was as low as $1331 MM dollars. The most profitable BTA refineries input hardwood, followed by switchgrass, and finally corn stover. A hardwood biorefinery that produces 2000 MT/day has a net present value of over $1200 MM. A similar scale switchgrass biorefinery has an NPV of over $750 MM, while a biorefinery inputting corn stover with a similar capacity has a NPV of over $550 MM dollars. The process synthesis framework described has the ability to produce any combination of xylene isomers and three sets of aromatics products were considered that placed no restriction on the type of xylenes produced, limited xylenes output to only the ortho-xylene isomer, and limited xylenes

This paper proposes novel methods of producing aromatics from biomass. A deterministic global optimization-based process synthesis framework was proposed to determine the most profitable process topology of a biomass to aromatics refinery. Several commercial, novel, and competing technologies were designed and described to convert biomass into aromatics via methanol. Three types of biomass (agricultural residues, perennial crops, and forest residues) were investigated to determine the effect that the type of biomass has on the overall profit of the refinery. Three sets of product compositions across five refinery scales were highlighted. Plant profitability, net present value, material and energy balances, and carbon and greenhouse gas balances were illustrated for all case studies investigated. W

DOI: 10.1021/acs.energyfuels.6b00619 Energy Fuels XXXX, XXX, XXX−XXX

X

biomass gasoline LPG benzene toluene p-xylene o-xylene

biomass gasoline LPG benzene toluene p-xylene o-xylene m-xylene vented CO2 seq CO2 LGHG GHGAF GHGAE GHGAC GHGI

biomass gasoline LPG benzene toluene p-xylene o-xylene m-xylene vented CO2 seq CO2 LGHG GHGAF GHGAE GHGAC GHGI

U-200

−12.26 1.54 0.00 0.61 0.00 5.25 0.00 0.00 6.43 0.00 1.57 2.10 −1.32 8.85 0.16

−6.14 0.77 0.00 0.31 0.00 2.62 0.00 0.00 3.22 0.00 0.79 1.05 −0.66 4.43 0.16

−21.10 1.54 0.00 0.61 0.00 5.25 0.00

U-200

U-100

U-100

−12.30 1.54 0.00 0.61 0.00 5.25 0.00 0.00 5.99 0.00 1.09 2.10 −1.28 8.85 0.11

−6.05 0.80 0.00 0.27 0.02 2.60 0.00 0.00 2.89 0.00 0.53 1.09 −0.64 4.37 0.11

−10.55 0.77 0.00 0.31 0.00 2.62 0.00

U-200

U-100

−52.78 3.85 0.00 1.53 0.00 13.11 0.00

U-500

−30.64 3.85 0.00 1.53 0.00 13.11 0.00 0.00 16.07 0.00 3.92 5.25 −3.29 22.13 0.16

U-500

−30.72 3.85 0.00 1.53 0.00 13.11 0.00 0.00 14.95 0.00 2.73 5.25 −3.19 22.13 0.11

U-500

−105.55 7.70 0.00 3.06 0.00 26.23 0.00

U-1000

−63.75 7.70 0.00 3.06 0.00 26.23 0.00 0.00 34.95 0.00 8.18 10.50 −5.09 44.27 0.16

U-1000

−60.45 8.02 0.00 2.67 0.24 25.97 0.00 0.00 28.89 0.00 5.34 10.95 −6.38 43.66 0.11

U-1000

−219.30 15.39 0.00 6.12 0.00 52.45 0.00

U-2000

−127.72 15.39 0.00 6.12 0.00 52.45 0.00 0.00 70.14 0.00 16.38 21.00 −10.21 88.54 0.16

U-2000

−121.03 16.05 0.00 5.36 0.00 52.40 0.00 0.00 57.92 0.00 10.69 21.89 −12.78 87.32 0.11

U-2000

−9.69 0.79 0.00 0.33 0.00 0.00 2.58

O-100

−5.60 0.78 0.00 0.32 0.00 0.00 2.59 0.00 2.61 0.00 0.71 1.07 −0.66 4.40 0.15

O-100

−5.62 0.78 0.00 0.33 0.00 0.00 2.59 0.00 2.42 0.00 0.49 1.07 −0.64 4.40 0.10

O-100

Table 14. Greenhouse Gas (GHG) Balances for the Case Studies Are Showna

O-500

−19.03 1.54 0.00 0.61 0.00 0.00 5.25

O-200

−51.37 3.86 0.00 1.51 0.00 0.00 13.12

O-500

−12.01 −44.92 1.54 3.86 0.00 0.00 0.60 1.51 0.00 0.00 0.00 0.00 5.25 13.12 0.00 0.00 6.15 32.27 0.00 0.00 1.54 5.84 2.10 5.26 −0.75 2.47 8.85 22.12 0.15 0.20 Switchgrass

O-200

−11.08 −29.19 1.54 3.86 0.00 0.00 0.60 1.51 0.00 0.00 0.00 0.00 5.25 13.12 0.00 0.00 4.66 13.29 0.00 0.00 0.98 2.58 2.10 5.26 −1.30 −2.23 8.85 22.12 0.10 0.10 Hardwood

O-500

Corn Stover O-200

−102.63 7.71 0.00 3.02 0.00 0.00 26.24

O-1000

−60.03 7.71 0.00 3.02 0.00 0.00 26.24 0.00 30.73 0.00 7.68 10.52 −3.74 44.23 0.15

O-1000

−58.62 8.02 0.00 2.66 0.23 0.00 25.99 0.00 26.88 0.00 5.17 10.95 −3.91 43.66 0.10

O-1000

O-2000

−205.53 15.43 0.00 6.05 0.00 0.00 52.48

O-2000

−120.11 15.43 0.00 6.05 0.00 0.00 52.48 0.00 61.52 0.00 15.36 21.05 −7.48 88.47 0.15

O-2000

−118.06 16.05 0.00 5.33 0.46 0.00 51.98 0.00 54.66 0.00 10.41 21.89 −7.35 87.32 0.10

−9.89 0.80 0.00 0.27 0.02 0.00 0.00

M-100

−6.04 0.80 0.00 0.37 0.00 0.00 0.00 2.52 3.12 0.00 0.77 1.09 −0.66 4.37 0.16

M-100

−5.76 0.80 0.00 0.27 0.02 0.00 0.00 2.60 2.57 0.00 0.51 1.09 −0.64 4.37 0.11

M-100

M-200

−19.77 1.60 0.00 0.53 0.05 0.00 0.00

M-200

−17.96 1.60 0.00 0.73 0.00 0.00 0.00 5.06 12.90 0.00 2.33 2.19 0.80 8.74 0.20

M-200

−11.51 1.60 0.00 0.53 0.05 0.00 0.00 5.20 5.15 0.00 1.01 2.19 −1.28 8.73 0.11

M-500

−55.11 4.01 0.00 1.33 0.12 0.00 0.00

M-500

−32.20 4.01 0.00 1.33 0.12 0.00 0.00 12.99 17.87 0.00 4.12 5.47 −1.15 21.83 0.16

M-500

−31.46 4.01 0.00 1.33 0.12 0.00 0.00 12.99 15.80 0.00 2.79 5.47 −1.44 21.83 0.11

−110.24 8.02 0.00 2.67 0.23 0.00 0.00

M-1000

−64.29 8.02 0.00 2.67 0.23 0.00 0.00 25.98 35.62 0.00 8.23 10.95 −2.29 43.66 0.16

M-1000

−63.65 8.02 0.00 2.67 0.23 0.00 0.00 25.98 32.39 0.00 5.64 10.95 −2.47 43.66 0.11

M-1000

M-2000

−220.58 16.05 0.00 5.33 0.47 0.00 0.00

M-2000

−130.84 15.41 0.00 6.07 0.00 0.00 0.00 52.47 73.69 0.00 16.80 21.03 −4.56 88.49 0.16

M-2000

−127.78 16.05 0.00 5.33 0.47 0.00 0.00 51.96 65.31 0.00 11.34 21.89 −4.55 87.32 0.11

Energy & Fuels Article

DOI: 10.1021/acs.energyfuels.6b00619 Energy Fuels XXXX, XXX, XXX−XXX

Article

The total GHG emissions (in CO2 equivalents - kg CO2 eq/s), as well as the GHG emissions avoided from liquids production (GHGAF), emissions due to electricity usage (GHGAE), GHG emissions avoided from chemicals production (GHGAC), and overall GHG emissions (LGHG), are also illustrated.

Table 15. NPVs for the Unrestricted 2000 MT/day Aromatics Refineries Cornstover U-2000 price ($/dry metric ton) NPV (MM $) Hardwood U-2000 price ($/dry metric ton) NPV (MM $) Switchgrass U-2000 price ($/dry metric ton) NPV (MM $)

100 786

120 552

140 310

60 1340

70 1221

80 1102

80 1023

100 778

120 532

output to only the meta-xylene isomer. The BTA refineries had superior life cycle greenhouse gas emissions compared to fossilfuel based processes.

■

ASSOCIATED CONTENT

S Supporting Information *

The Supporting Information is available free of charge on the ACS Publications website at DOI: 10.1021/acs.energyfuels.6b00619. Appendix A and B (PDF)

■

AUTHOR INFORMATION

Corresponding Author

*Tel: (979) 458-0253. E-mail: fl[email protected]. Notes

The authors declare no competing financial interest.

■

ACKNOWLEDGMENTS The authors acknowledge partial financial support from the National Science Foundation (NSF EFRI-0937706 and NSF CBET-1158849).

■

REFERENCES

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a

25.98 32.57 0.00 −40.77 10.95 −2.07 43.66 −0.78 12.99 16.28 0.00 −20.38 5.47 −1.03 21.83 −0.78

51.96 65.20 0.00 −81.56 21.89 −4.13 87.32 −0.78

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Switchgrass

O-200 O-100 U-2000 U-1000 U-500 U-200 U-100

Table 14. continued

m-xylene vented CO2 seq CO2 LGHG GHGAF GHGAE GHGAC GHGI

M-100

M-200

M-2000

Energy & Fuels

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