Coal and Biomass to Liquid Transportation Fuels - American Chemical

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Coal and Biomass to Liquid Transportation Fuels: Process Synthesis and Global Optimization Strategies Alexander M. Niziolek,† Onur Onel,† Josephine A. Elia,† Richard C. Baliban,† Xin Xiao,*,‡ and Christodoulos A. Floudas*,† †

Department of Chemical and Biological Engineering, Princeton University, Princeton, New Jersey 08544, United States Langfang Engineering and Technology Centre, National Engineering Laboratory for Hydrometallurgical Cleaner Production Technology, Institute of Process Engineering, Chinese Academy of Sciences, Beijing, 100190, China

‡

S Supporting Information *

ABSTRACT: The thermochemical conversion of coal and biomass to liquid transportation fuels from a synthesis gas intermediate is investigated using an optimization-based process synthesis framework. Two distinct types of coal (LV bituminous and coal commonly found in the province of Anhui, China) and two types of biomass (hardwood and duckweed) are considered as feedstocks. The superstructure incorporates alternative conversion pathways of synthesis gas which include methanol formation and conversion into Fischer−Tropsch hydrocarbons. Methanol may be converted to gasoline or olefins, and the olefins may be subsequently converted to gasoline and distillate. A rigorous deterministic global optimization branch-and-bound framework is utilized to determine the optimal process topology that produces liquid fuels at the lowest possible cost. Economies of scale are evident as the refinery capacity increases and it is observed that the fuel ratios of the final liquid products have a significant impact on the optimal topology of the plant. The results suggest that liquid fuels production from coal and biomass can be competitive with petroleum-based processes.

1. INTRODUCTION The United States faces major challenges due to the high energy demand that its residential, commercial, industrial, and transportation sectors require. In 2012, the U.S. consumed 18,490 thousand barrels of petroleum products per day (kBD).1 The U.S. transportation sector accounted for the majority of this consumption (13,014 kBD) and is projected to continue to do so until 2040.1,2 Current projections predict a continuous increase in crude oil production in the U.S. until 2020 due to high oil prices and technological advances in production from shale and tight oil formations.2 However, these outlooks are highly uncertain due to the early stages of tight oil development.2 In addition, the instability of global oil prices, political unrest in the Middle East, and a growing concern over greenhouse gas emissions compound the problems facing the United States and, in particular, the U.S. transportation sector. These challenges spur the development of processes that produce liquid transportation fuels from domestically available carbon-based feedstocks. Coal, biomass, and natural gas are three major feedstocks that the United States has at its disposal to gradually replace part of petroleum as its primary energy source. Coal is an attractive precursor to liquid fuels because its delivered cost ($2.0−$2.5/ MM Btu) is less expensive than that of biomass ($4.0−$9.0/ MM Btu) or natural gas ($4.8−$5.8/MM Btu). 3 The tremendous interest in utilizing biomass as a feedstock stems from the fact that it is a renewable carbon resource that has the ability to reduce greenhouse gas emissions through the capture of CO2 during photosynthesis.4−6 Recent reviews have highlighted the technologies of producing first- and second-generation biofuels.7,8 Recent advances in the hydraulic fracturing and horizontal drilling of natural gas, together with the discovery of © XXXX American Chemical Society

new sources of shale gas in the U.S., have increased interest in this feedstock as a precursor to liquid fuels. A recent review has highlighted the key contributions to the production of liquid transportation fuels from single and hybrid feedstock energy systems utilizing coal, biomass, and natural gas.9 Hybrid energy systems that have been investigated in the literature include coal and natural gas to liquids (CGTL),10−18 biomass and natural gas to liquids (BGTL),19−23 and coal, biomass, and natural gas to liquids (CBGTL).24−29 Coal and biomass to liquid transportation fuels (CBTL) is the type of hybrid energy feedstock considered in this paper. The higher carbon content in coal results in a higher conversion to CO2, which, taken over the entire life-cycle of a conventional coal-toliquids (CTL) refinery, amounts to almost twice the life-cycle CO2 emissions of a crude-oil based plant.30−33 Since this is one prevailing detriment to widespread development of coal-toliquids plants, it is important to investigate hybrid systems of coal and biomass that can take advantage of the economic and environmental benefits that both of these feedstocks can provide. Two distinct types of biomass are considered in this paper: hardwood and duckweed. Duckweed is a fast-growing, floating aquatic plant that grows in wastewater streams.34−48 Much of the focus in literature has been on the conversion of duckweed to ethanol;41−44,47 while some studies have studied pyrolysis37−39,49 and hydrothermal processing36 of duckweed. Recently, Baliban et al. Special Issue: Jaime Cerdá Festschrift Received: February 4, 2014 Revised: April 11, 2014 Accepted: April 24, 2014

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dx.doi.org/10.1021/ie500505h | Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

Industrial & Engineering Chemistry Research

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methanol-to-gasoline (MTG) or methanol-to-olefins (MTO), (v) hydrocarbon upgrading via ZSM-5 zeolite catalysis, olefin oligomerization, or carbon number fractionation and subsequent treatment. Gasoline, diesel, and kerosene will be the major products from the refinery. Liquefied petroleum gas (LPG) and electricity may be sold as byproducts.

investigated the thermochemical conversion of duckweed into liquid transportation fuels.48 In addition, two distinct types of coal are compared in this paper: LV bituminous, as well as a coal commonly found in the province of Anhui, China. Since duckweed is available near this province it is important to investigate the viability of CBTL systems across multiple geographies where the feedstock compositions are fairly different. Coal and biomass to liquid fuels processes have previously been investigated in the literature. Warren and El-Halwagi investigated the production of liquid fuels and hydrogen from coal and municipal solid waste.50 Kreutz et al. proposed cofeeding coal with biomass and producing liquid fuels through Fischer−Tropsch refining.51 Liu et al. expanded upon this work to include process simulations and economic analysis for coal and/or biomass hybrid systems.52 Larson et al. investigated the production of liquid fuels and electricity from coal and biomass.53,54 While the authors thoroughly explained the conceptual design of their CBTL plants, they only considered a fixed process topology, simulated the process to obtain mass and energy balances, and then completed the economic analysis. Chen et al. proposed a nonlinear programming model that maximized the net present value of a polygeneration system that could produce power, liquid fuels, and chemicals from coal and biomass.55,56 However, Chen et al. only considered liquid fuels production from Fischer− Tropsch refining. It is important to consider multiple conversion technologies that can produce liquid transportation fuels. The development of process synthesis strategies that can efficiently determine the optimal process topology from thousands of scenarios have now made this approach possible.26−29 Therefore, this study focuses on the optimization-based framework for the production of liquid transportation fuels from coal and biomass. A large-scale nonconvex mixed integer nonlinear optimization (MINLP) model is utilized to analyze the environmental and economic trade-offs of the CBTL process. The process superstructure includes several alternatives that have been described in detail in previous works;23−29,48,57−59 however, all key components of the CBTL refinery will be discussed in the following section. Simultaneous heat, power, and water integration25,26,60−64 is included that utilizes heat engines to convert waste heat into electricity and treats and recycles wastewater. Each unit in the CBTL refinery is rigorously modeled to ensure proper operation. Additionally, the electricity and utility requirements for each unit, as well as for the mechanical pretreatment of the feedstocks, are included in the process superstructure. The process synthesis framework proposed uses a global optimization branch-and-bound strategy to determine the optimal plant design.28 The value of the objective function of the process topology selected is mathematically guaranteed to be within a small percentage of the best possible value. This paper is the first that investigates the process synthesis of coal and biomass systems utilizing multiple synthesis gas conversion technologies with a global optimization framework. Previous studies that fix the process topology are difficult to compare and analyze. This study presents a baseline that can provide a fair comparison between different process alternatives and mathematically guarantees that the topology design selected is within a few percentage points of best possible value. The process synthesis framework for the CBTL process will include (i) biomass gasification with/without recycle synthesis gas, (ii) coal gasification with/without recycle synthesis gas, (iii) synthesis gas conversion via Fischer−Tropsch (FT) synthesis or methanol synthesis, (iv) methanol conversion via

2. CBTL PROCESS SUPERSTRUCTURE: CONCEPTUAL DESIGN AND MATHEMATICAL MODELING Baliban et al. and Elia et al. described each section of a process refinery that inputs either single or hybrid feedstocks of coal, biomass, and natural gas in previous works.23−29,48,57−59 The following subsections will detail the key components in the conceptual design of the CBTL refinery. 2.1. Biomass Handling and Gasification. The coal and biomass are input into separate gasification units, and only one type of biomass is input into the refinery in order to reduce the complexity of the gasification section. The CBTL refinery is designed to input hardwood biomass or duckweed with a representative composition that is obtained from the ECN Phyllis database and shown in Table 1. The hardwood is assumed Table 1. Feedstock Proximate and Ultimate Analysis for Hardwood and Duckweed66 proximate analysis (db, weight %)

heating values (kJ/kg)

feed type

moisture (ar)

hardwood duckweed

45 −

feed type

C

H

N

Cl

S

O

hardwood duckweed

50.19 38.93

5.9 595

0.32 5.78

0 1.32

0.03 0.99

41.42 38.38

HHVc

LHVd

2.14 N/A N/A 19 130 17.6 75.3 7.1 15 157 Ultimate Analysis (db, weight %)

17 842 13 972

ash

VMa

FCb

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

to be delivered as woodchips to the refinery and thus is subject to an initial screening to prevent sizes larger than 2 in. to be sent to the biomass gasifier.65 The larger particles are sent to a grinder for further size reduction.65 The duckweed is delivered to the refinery and air-dried for 2 days in order to reduce the moisture content (to 7.9 wt %) to an acceptable level prior to being fed into the gasifier.48 The biomass gasification flowsheet is shown in Figure 1. Prior to entry into the gasifier, the hardwood is first dried, and its moisture content is reduced to 20 wt %. Flue gas generated in the refinery is sent to the drier to supply the heat necessary for the drying. The flue gas exits the drier at 110 °C and 1.05 bar, passes through an air cyclone and baghouse filter to remove any particulates, and is vented.65 The dried biomass exits the drier at 105 °C and 1.05 bar and is transferred into the biomass gasifier via a lockhopper and compressed CO2 (10 wt %). Since the moisture content of the duckweed is less than 20 wt %, it can be directly lockhopped and transferred to the gasifier. The biomass gasifier operates at 30 bar and can input either solid biomass or a combination of biomass and recycle gases as fuel. The effluent from the biomass gasifier contains a mixture of synthesis gas (CO and H2), C1 and C2 hydrocarbons, ash, tar, char, and acid gases such as NH3, H2S.24,26 With the aid of cyclones, any solid ash and char from the gasifier effluent is separated from the vapor phase products and recycled back to the gasifier. The recycle of char B



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Figure 1. Biomass gasification flow diagram. If necessary, the biomass will first be dried to 20 wt % moisture and then transferred to the gasifier system using a lockhopper and compressed CO2. The gasifiers will input either a solid biomass or a combination of solid biomass and recycle gases as fuel. Solid ash and char that are formed in the gasifiers are separated from the vapor phase effluent using cyclones and recycled back into the gasifiers. The raw syngas is further reformed with a catalytic tar cracker.

tars, (iii) 99% of C2H6, (iv) 90% of C2H4, and (v) 90% of NH3.65 A pilot-scale demonstration of the unit is being installed to validate the performance over an extended period of time.67 Additional steam input into the tar cracker is not necessary because the steam present in the effluent from the gasifier is enough to reform the syngas.67 The circulating catalyst between the tar reformer reactor and the catalyst regenerator provides the heat for tar conversion.67 In addition to providing heat for the tar reforming reactions, the catalyst regenerator also restores catalyst activity by burning the coke off the catalyst; however, the heat generated from coke formation is insufficient for the endothermic reforming reactions, so supplemental combustion gases are sent through the regenerator.67 The raw syngas effluent from the tar cracker is directed to the syngas cleaning section. 2.2. Coal Handling and Gasification. The CBTL refinery is designed to input one type of coal to reduce the complexity of the handling and processing section. Two distinct types (LV bituminous, Anhui Province Coal) of coal are considered in this paper. LV bituminous coal has a representative composition that is obtained from the NETL Detailed Coal Specifications Report and is shown in Table 2, along with the specifications of the coal commonly found in Anhui, China. The delivered coal is sent into receiving hoppers and then fed directly into a vibratory feeder which reduces the size of the coal to 3 in.68 The coal is transferred onto a belt conveyor, is passed under a magnetic plate to remove tramp iron, and is then further crushed to reduce the size to 1.25 in.68 The coal handling and gasification flow diagram is illustrated in Figure 2. The coal from the province of Anhui, China, is sent to the drier to reduce its moisture content to the acceptable threshold of 2 wt %. The heat necessary for the drying is supplied by flue gas generated within the CBTL refinery. Since the moisture content of LV bituminous coal (0.65 wt %) is less than the maximum allowable threshold (2 wt %) for entry into the gasifier, the drying step is bypassed. Using a lockhopper and

assumes that effectively 100% of the carbon present in the biomass feedstock is converted to vapor products, while the ash is output as slag. The composition of the gasifier effluent is determined using the biomass gasifier model developed by Baliban et al. and is a function of the biomass gasifier temperature, oxidizer flow rate, and biomass composition.24,26 The gasifier will operate at either 900 °C, 1000 °C, or 1100 °C and will input steam to gasify the fuel (biomass or biomass/recycle gas) and reform the C1−C2 hydrocarbons and tar species. The gasifier is modeled using discrete temperatures in order to avoid introducing more nonlinearities into the process synthesis model. More specifically, the equilibrium constant of the water-gas-shift (WGS) reaction would need to be calculated using a nonlinear exponential function if the temperature range is continuous. However, the authors note that any number of discrete operating temperatures for the gasifier may be introduced. Additionally, every unit with WGS equilibrium is modeled using discrete temperatures. The heat needed for the reforming reactions will be provided by high-purity oxygen, which will additionally facilitate the cracking of the tar species. Operating at such high temperatures will help the synthesis gas effluent approach the water-gas-shift equilibrium, although the hydrocarbon concentrations in the effluent will be above their equilibrium values. Any CO2 present within the gasifier may be consumed by reacting with the H2 present within the gasifier through the reverse WGS reaction. Therefore, by channeling a portion of the H2 produced from pressure-swing adsorption or electrolysis of water, the CO2 produced in the process may be recycled into the gasifier.26 After treatment through the biomass ash cyclones, the ash-free effluent is sent to a catalytic tar cracker operating at 925 °C which reforms the (i) tar into H2 and CO, (ii) NH3 into N2 and H2, and (iii) C1 and C2 hydrocarbons into H2 and CO. The National Renewable Energy Laboratory’s (NREL) bench-scale performance of a tar cracker converts (i) 80% of the CH4, (ii) 99.6% of C



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water-gas-shift reactor prior to entering the scrubbing system or may be directly sent to the scrubbing system after being cooled to 185 °C. The water-gas-shift reactor operates at 26 bar, at either 300 °C, 400 °C, 500 °C, or 600 °C, and can facilitate either the forward water-gas-shift reaction to increase the H2/CO syngas ratio or the reverse water-gas-shift reaction to decrease the CO2 concentration. The forward water-gas-shift reaction will generate additional CO2 that will be subsequently removed in the syngas treatment units, while the exothermic heat of this reaction will be removed via steam generation. The operation of the reverse water-gas-shift reaction will require a high inlet concentration of CO2 (>0.5) that can be achieved through recycle of the CO2 from the process. Additionally, the pressure swing adsorption unit, electrolyzer, or autothermal reactor will provide the H2 necessary for the reverse water-gas-shift reaction. The electrolyzer or air separation unit will supply the oxygen necessary to provide heat for the reverse water-gas-shift reaction. The decision to include the dedicated water-gasshift reactor depends on the H2/(CO + CO2) ratio needed for syngas conversion, whether the necessary amount of H2 can be economically removed using pressure swing adsorption, whether the incorporation of an electrolyzer for H2 production is economical, and the specified lifecycle greenhouse gas emissions of the refinery. If it is selected, the effluent from the watergas-shift reactor is cooled to 185 °C and sent to the scrubbing section. The scrubbing system will remove any residual tar, particulates, and NH3 from the raw syngas. Any wastewater generated in the scrubber is treated in a biological digestor. The effluent from the scrubbing system is sent to an acid gas removal system that co-removes H2S and CO2 from the syngas. The acid gas removal system is a dual-capture methanol absorption system (Rectisol unit) which removes 100% of the H2S and 90% of the CO2 present in the input stream.51 The dual-capture system captures the H2S and CO2 in separate streams. The sulfur-rich stream will contain 3 mol of CO2 for every mole of H2S.51 The Rectisol system must be utilized to prevent catalyst poisoning in the subsequent hydrocarbon production/upgrading units due to the presence of sulfur in the syngas. The sulfur-rich stream from the dual-capture system is directed to the Claus recovery system.68

Table 2. Feedstock Proximate and Ultimate Analysis for the LV Bituminous and Anhui, China, Coal69 proximate analysis (db, weight %) feed type

moist. (ar)

LV bituminous Anhui, China

0.65 4.08

HHVc

LHVd

4.77 19.27 75.96 34 946 18.34 43.32 45.55 25 300 ultimate analysis (db, weight %)

34 040 −

ash

VMa

FCb

heating values (kJ/kg)

feed type

C

H

N

Cl

S

O

LV bituminous Anhui, China

86.71 62.54

4.23 4.33

1.27 1.12

0.19 0

0.66 0.43

2.17 10.03

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

compressed CO2, the coal is then transferred to the high-pressure (31 bar) coal gasifier. Similar to the operation of the biomass gasifier, the coal gasifier can input either solid coal or a combination of solid coal and recycle gases as fuel. The effluent of the coal gasifier is determined by using the stoichiometric coal gasifier model proposed by Baliban et al. which assumes that the effluent is in equilibrium with respect to the water-gas-shift reaction.24,26 The syngas effluent composition is a function of the coal gasifier temperature, oxidizer flow rate, and coal composition.24,26 The coal gasifier will operate at either 1100 °C, 1200 °C, or 1300 °C and will input steam or oxygen to gasify the fuel (coal or coal/recycle gas). The heat needed for the reforming reactions will be provided by high-purity oxygen. Analogous to the operation of the biomass gasifier, a portion of the H2 produced from pressure-swing adsorption or electrolysis of water may be sent to the solid-/vapor-fueled coal gasifier, allowing the CO2 produced in the process to be recycled into the coal gasifier. The syngas effluent is quenched and then sent to the coal ash cyclones. The ash and char are separated from the vapor-phase products and recycled back into the gasifier, where the ash is ultimately output as slag. The ash-free syngas is directed to the cleaning section. 2.3. Synthesis Gas Cleaning. As shown in Figure 3, the coal and biomass syngas may be sent to a dedicated sour

Figure 2. Coal gasification flow diagram. The coal from Anhui, China, is sent to the drier, while the drying step is bypassed for LV bituminous coal. The coal is then transferred to the gasifier via a lockhopper. The gasifiers will input either solid coal or a combination of solid coal and recycle gases as fuel. The syngas effluent is quenched, and any ash or char that formed during gasification is separated using the coal cyclones and recycled back. The raw syngas is then split to either the syngas scrubber or to a dedicated reverse water-gas-shift reactor. D



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Figure 3. Synthesis gas (syngas) cleaning flowsheet. Syngas can be passed over either a forward or reverse water-gas-shift reactor to alter the H2 to CO/CO2 ratio prior to the Hydrocarbon Production section. The syngas is then cooled down and directed to a dual-capture methanol-based unit in order to remove the CO2 and H2S in the raw syngas. The H2S gases are then directed to a Claus plant for recovery of H2S as elemental sulfur. The captured CO2 may be either vented, recycled back into the refinery, or sequestered.

Iron-based Fischer−Tropsch reactors will facilitate the watergas-shift reaction and could consume recycled CO2 by inputting H2 produced in the refinery to produce the CO needed for the FT reaction.26,27,29 The consumption of CO2 will occur if the ratio of CO2/(CO+CO2) is above some critical threshold; otherwise, CO will likely be converted to CO2 within the units.48,59 The critical threshold depends on the relative amount of H2 with respect to CO and CO2. Iron-based LTFT units have been successfully operated with inlet H2/CO ratios of 0.5−170−72 and produce effluents with a H2/CO ratio between 1.7 and 2.0. However, as much as 50% of the CO is converted to CO2. By setting the H2/(CO+CO2) ratio such that CO2 is employed as a carbon source in the reverse water-gas-shift reaction, the CO2 concentration in the effluent can be controlled.70 The Ribblett ratio70,73 is defined such that the ratio of H2/(2CO +3CO2) is approximately 1 and is highly beneficial because the unreacted syngas in the effluent of the FT reactor will have the same composition as the inlet. Therefore, internal and external gas loops may be designed in the CBTL refinery that allow for high conversion rates of CO and CO2. The H2/(CO+CO2) ratio will be examined in the following two ways. Two iron-based FT units will operate at either low temperature (240 °C) (LTFTRGS) or high temperature (320 °C) (HTFTRGS) and will facilitate the reverse watergas-shift reaction by requiring an inlet Ribblett ratio of 1. The other two iron-based FT (medium-temperature nominal wax FT, MTFTWGS-N, and medium-temperature minimal wax FT, MTFTWGS-M) units will operate at 267 °C and require an inlet H2/CO ratio between 0.5 and 0.7. The forward water-gas-shift reaction will ensure that the effluent H2/CO ratio is approximately 1.7. The effluent composition from these two units is based off of units developed by Mobil Research and Development.29,71,72 Hydrogen can be sent to any one of the cobaltbased or iron-based reactors to shift the H2/CO or H2/CO2 ratios. The two iron-based FT reactors (MTFTWGS-N and

The Claus recovery system converts 95% of the H2S and SO2 into solid sulfur.68 The tail gas from the Claus recovery system will be hydrogenated to H2S and recycled back to the acid gas removal systems. Thus, 100% recovery of the sulfur in the CBTL process can be achieved. The CO2 from the acid gas recovery units may be either (i) compressed to 31 bar and recycled in the process, (ii) compressed to 150 bar and sequestered, or (iii) vented to the atmosphere. Interstage cooling will be utilized to ensure that the temperature of each compressor outlet remains below 200 °C. 2.4. Hydrocarbon Production and Upgrading. 2.4.1. Fischer−Tropsch Hydrocarbon Production. The clean syngas is split to either the Fischer−Tropsch (FT) reactors (as shown in Figure 4) or to the methanol synthesis reactor. Six different Fischer−Tropsch reactors, each operating at a pressure of 20 bar, are considered in the process synthesis superstructure. Cobalt-based and iron-based FT reactors will both be considered and are discussed in this section. High per-pass conversions of CO to Fischer−Tropsch liquids are attainable in the Co-based FT reactors since they do not facilitate the water-gas-shift reaction. However, deactivation of Co-based FT catalysts, which has been attributed to carbon disposition and fouling, oxidation and mixed-oxide formation, and changes in Co crystallite size, poses a serious problem because it alters product selectivity.70 Carbon deposition prevents access to catalytic sites and is characterized by increased methane selectivity; however, it can be reversed through rejuvenation or regeneration of the catalyst.70 The deactivation of Co-catalysts due to oxidation is a largely debated topic, and its effect on catalyst stability is unclear.70 Thus, to avoid catalyst oxidation, the CO per-pass conversion is set to 60% in this study, but per-pass conversions of 80% can be achieved if catalyst stability can be attained.70 Only the cobalt low-temperature FT reactor (LTFT) has been commercially available to date, but the high-temperature unit is also considered in the superstructure with a typical alpha value of 0.72 that is consistent with high-temperature FT (HTFT) operation. E



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Figure 4. Fischer−Tropsch hydrocarbon production flow diagram. Clean syngas may be sent to either the high-temperature cobalt-based FT (HTFT), low-temperature cobalt-based FT (LTFT), high-temperature iron-based FT (HTFTRGS), low-temperature iron-based FT (LTFTRGS), mid-temperature iron-based nominal wax FT (MTFTWGS-N), or mid-temperature iron-based minimal wax FT (MTFTWGS-M) reactors. The latter two units will facilitate the forward water-gas-shift reaction. The vapor-phase effluents are sent to be upgraded, while the wax is sent to a hydrocracker.

into gasoline-range hydrocarbons and some distillate.71,72 No additional treatment of the oxygenates is necessary since the ZSM-5 unit will be able to convert these compounds into additional hydrocarbons. The effluent from the ZSM-5 reactor is modeled after one previously presented by Mobil.71,72 The raw effluent from the ZSM-5 reactor is then sent to a fractionation unit (Figure 5) that separates the water and distillate from the gasoline product. The sour water is sent to the wastewater treatment section, the distillate is sent to be hydrotreated, and the gasoline is sent to be processed in the LPG−gasoline separation section, described later in the text. 2.4.3. Methanol Synthesis. Alternatively, the clean syngas may be compressed to 51 bar and directed to the methanol synthesis reactor, as shown in Figure 7. The reactor operates at 300 °C and 50 bar and will input synthesis gas with a Ribblett ratio equal to 1. The reactor assumes that equilibrium exists between the water-gas-shift reaction (eq 1) and the methanol synthesis reaction (eq 2):65

MTFTWGS-M) that facilitate the forward water-gas-shift reaction can additionally input steam to shift the H2/CO ratio within the reactor. The FT reactors will output two streams: a liquid-phase effluent containing wax which is sent to a hydrocracker and a vapor-phase effluent containing C1−C30+ hydrocarbons that is upgraded, as described in the following section. 2.4.2. Fischer−Tropsch Hydrocarbon Upgrading. As illustrated in Figure 5, the vapor-phase FT effluent can be either sent to a series of treatment units to knock out the water oxygenates present in the stream or passed over a ZSM-5 catalytic reactor. In the former case, the water-soluble oxygenates are first stripped from the stream and then sent to a three-phase separator to separate the aqueous phase from the residual vapor and hydrocarbon liquid. The vapor phase is then sent to a third unit to remove any additional oxygenates present in the stream. The oxygenates and water are sent to the wastewater treatment section. The water-lean FT hydrocarbons are then directed to a hydrocarbon recovery system (shown in Figure 6) where they are split into C3−C5 gases, naphtha, kerosene, distillate, wax, offgas, and wastewater.24,74 Each stream will be upgraded following the Bechtel design.74,75 The upgrading units include a wax hydrocracker, a distillate hydrotreater, a kerosene hydrotreater, a naphtha hydrotreater, a naphtha reformer, a C4 isomerizer, a C5/C6 isomerizer, a C3/C4/C5 alkylation unit, and a saturated gas plant. Baliban et al. described these upgrading units and their corresponding effluents in more detail in previous works.24,29 The ZSM-5 catalytic reactor operates at a temperature of 408 °C and a pressure of 16 bar71 and will convert the FT effluent

CO + H 2O ↔ CO2 + H 2

(1)

CO + 2·H 2 ↔ CH3OH

(2)

The amount of water generated in the output stream of the methanol synthesis reactor depends heavily on the input concentration of CO2. Additional water will be generated through the reverse water-gas-shift reaction if the input CO2 increases. In addition, the per-pass conversion of the CO + CO2 would decrease. This can be reconciled if the input stream has a Ribblett ratio of 1. The CO and CO2 could then be recycled back to the reactor in order achieve a per-pass conversion of F



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Figure 5. Fischer−Tropsch hydrocarbon upgrading flow diagram. The Fischer−Tropsch vapor effluent may be passed over a series of treatment units to remove the water-soluble oxygenates or passed over a ZSM-5 catalytic reactor to produce gasoline and distillate.

Figure 6. Fischer−Tropsch hydrocarbon-upgrading flow diagram. The water-lean Fischer−Tropsch hydrocarbons are sent to a series of upgrading units which separate and upgrade the hydrocarbons into distillate and gasoline-range products.

so large amounts of water in the input are not assumed to be an issue.76 The raw methanol effluent leaving the reactor is first cooled to 35 °C and then sent to a flash unit to remove a majority (95%) of the methanol using vapor−liquid equilibrium. The vapor phase is

CO + CO2 greater than 90%. The methanol conversion units that will input the crude methanol are assumed to be able to handle as much as 50 wt % water. Thus, no further purification units for the methanol are necessary. The methanol processing units will produce roughly 50 wt % water from the hydrocarbon synthesis, G



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Figure 7. Methanol synthesis and conversion flow diagram. The clean syngas is sent to a methanol synthesis reactor. The crude methanol can then be split to either the methanol-to-gasoline (MTG) process or the methanol-to-olefins (MTO) and Mobil olefins-to-gasoline/distillate (MOGD) processes.

kerosene. Prior to entering the MTO unit, the methanol is heated to 375 °C at 1.2 bar. The MTO unit operates at a temperature of 375 °C and 1 bar.76,80 At such high temperatures, the formation of light olefins is thermodynamically favored.76 The MTO reactor uses a SAPO-34 catalyst that has a high selectively to light olefins while achieving complete conversion of methanol.80 The MTO effluent contains over 95 wt % C2−C4 olefins80 and is sent to be fractionated (MTO-F) so that 100% of the olefins are sent to the MOGD unit; 100% of the C1−C3 hydrocarbons are recycled back to the refinery, 100% of the C5+ species are blended with the gasoline pool, and 100% of the water generated is sent for treatment. The MOGD unit is a fixed bed reactor that can operate at different modes and is able to convert the olefins to gasoline and distillate.76,81 The two modes described in literature for the MOGD process produced up to 82 wt % of distillate in the max distillate mode and up to 84 wt % of gasoline in the max gasoline mode.76 By shifting the reaction temperature and recycle composition, this unit can be used to maximize the production of diesel or gasoline. The gasoline and diesel selectivity is greater than 95% of the olefin feed.76 The MOGD reactor operates in the presence of a ZSM-5 catalyst and typically operates at 400 °C and 1 bar. It has also been shown that at constant temperature and pressure, the product distribution for different olefin feeds is identical.81 Low pressure steam is generated in order to handle the exothermic heat of reaction.58 An atom balance is used to model this unit with an allowable gasoline/distillate product ratio that ranges from 0.12 to >100.76 The MOGD effluent will be fractionated to separate the diesel and kerosene cuts from the gasoline and light gases; 100% of the C11−C13 species are separated as kerosene and 100% of the C14+ species are separated as diesel. 2.4.5. LPG−Gasoline Separation. The process flow diagram for the separation of the LPG and gasoline generated from either

split so that 95% of the stream is recycled back to the methanol synthesis reactor, while the remaining 5% is used as fuel gas in the process. The crude methanol from the flash unit is heated to 300 °C and passed over a turbine, where the gas is expanded to 5 bar in order to recover electricity. The turbine effluent is cooled to 60 °C and directed to a degasser distillation column which will remove all of the entrained gases and recover 99.9 wt % of the methanol. The entrained gases will be recycled and used as fuel gas in the process. The bottoms product leaving the degasser will be a mixture of methanol and water. 2.4.4. Methanol Conversion. The methanol produced from the reactor is split to either the methanol-to-gasoline (MTG) process or the methanol-to-olefins (MTO) and Mobil olefins-togasoline/distillate (MOGD) processes.65,76−79 The MTG process will utilize a ZSM-5 zeolite to catalytically convert the methanol to gasoline range hydrocarbons in a fluidized bed reactor. The methanol input will first be raised to 14.5 bar and preheated to 330 °C before entering the adiabatic reactor that operates at 400 °C and 12.8 bar. The MTG effluent is described in Table 3.4.2 of a Mobil study77 and in Process Flow Diagram P850-A1402 of the National Renewable Energy Laboratory (NREL) study.65 The MTG reactor is modeled with an atomic balance and an effluent composition equivalent to the one presented in the NREL study. All of the methanol entering the MTG reactor is converted, and the effluent contains 56 wt % water and 44 wt % hydrocarbons.65,77 The crude hydrocarbon product is composed of 2 wt % light gases, 19 wt % C3−C4 gases, and 79 wt % gasoline product and will be sent to be upgraded in the LPG−gasoline separation section. The final products will contain 82 wt % gasoline and 10 wt % LPG, with the balance being recycle gases.65 Alternatively, the methanol may be split to the methanol-toolefins reactor and subsequently to the Mobil olefins-to-gasoline/ distillate process to produce a mixture of gasoline, diesel, and/or H



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Figure 8. LPG−gasoline separation flowsheet.

The external gas loop configuration, shown in Figure 9, consists of three processing options that include an autothermal reformer, a fuel combustor, and a gas turbine. The autothermal reformer will convert the C1−C2 hydrocarbons into syngas that can be recycled back to the FT or methanol synthesis reactors. Alternatively, the light gases may be directed to a fuel combustor to provide heat or to a gas turbine to provide electricity for the process units. The effluent from the fuel combustor or the gas turbine is cooled to 35 °C, directed to a water knockout unit, and either passed through a CO2 recovery unit or vented to the atmosphere. 2.6. Hydrogen/Oxygen Production. As shown in Figure 10, high-purity hydrogen is supplied by pressure-swing adsorption or electrolysis of water. Additionally, oxygen will be provided by an air separation unit or electrolysis of water. 2.7. Wastewater Treatment. The complete wastewater treatment network for the CBTL refinery is illustrated in Figures 11 and 12. Process wastewater is directed to either a biological digestor or sour stripper to remove the H2S, NH3, and oxygenates in the wastewater streams. The sulfur-rich streams are directed to the Claus plant to recover and remove sulfur, while additionally providing heat for steam production. The output from the wastewater network includes outlet wastewater that meets all specifications for discharge,27 steam to process units in the refinery, and process water to the electrolyzers. 2.8. Unit Costs. Each unit in the CBTL refinery has a total direct cost, TDC, which is estimated using cost parameters from several literature sources.51,65,68,71,72,74,77,82−85 The tabulated values of the cost parameters are outlined in previous studies by Baliban et al.23,24,26−29,48,57−59 and are utilized in eq 3:

the ZSM-5 conversion of the Fischer−Tropsch hydrocarbons or the methanol conversion units is shown in Figure 8. Initially, light gases are first removed in one of two knockout units, and the flashed liquid is directed to a de-ethanizer that removes the light hydrocarbons from the gasoline. The light gases are sent to an absorber column that utilizes lean oil from a downstream unit as the absorbing liquid.65 The light gases from the absorber are recycled as fuel gas in the refinery, while the lean oil and absorbed hydrocarbons are recycled back to the de-ethanizer as reflux.65 The bottoms from the de-ethanizer are sent to a stabilizer column that removes the butanes from the gasoline. The butanes are directed to an alkylation unit that produces iso-octane. It was assumed that the alkylate was iso-butane, and that butene was completely converted to iso-butane.65 The effluent from the alkylation unit is separated into LPG and gasoline. The bottoms from the stabilizer column are directed to a splitter column that provides a portion of the lean oil for the absorbing column. Utility requirements for these units are obtained from the NREL report.65 2.5. Light Gas Handling. Light gases from the CBTL refinery consist of unreacted syngas from either the hydrocarbon production section or fuel gases in the saturated gas plant effluent. The light recycle gases are composed of C1−C2 hydrocarbons, unreacted syngas, and inert species. Because the light gases exiting the Fischer−Tropsch or methanol synthesis units may contain large amounts of H2 and CO, they may be recycled back to the synthesis units in an internal gas loop configuration that increases the per-pass conversion to liquid fuels. To prevent the buildup of inert species in this configuration, a portion of the light gases must be purged. The light gases that are purged, or the gases that are not directly recycled back in the internal gas loop configuration, may be processed in an external gas loop configuration before potentially being sent back to the synthesis units.

TDC = (1 + BOP) ·Co· I

Srsf So

(3)



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Figure 9. Light gas handling flowsheet.

Figure 10. Hydrogen and oxygen production section. The hydrogen and oxygen are first passed through heat exchangers (denoted by X) before being directed to process units in the CBTL refinery. J



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Figure 11. Process wastewater treatment section.

Figure 12. Utility wastewater treatment section.

scaling cost factor. The Chemical Engineering Plant Cost Index is used to convert the costs into 2012 dollars.86 To calculate the total plant costs, the total direct costs (TDC) are added with the indirect costs (ICs), which include

where the balance of plant (BOP) percentage (site preparation, civil works, etc.) is assumed to be 20% of the total installed unit 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 K



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reaction was equal to 16, while units restricted by the steam reforming reactions or methanol synthesis reactions were partitioned into 8 pieces. All of the splitter units are partitioned using 8 segments. Binary variables are introduced to activate only one segment of the domain. The reader is directed to a previous study by Baliban et al. for a more detailed explanation of the underestimation of the nonconvex nonlinearities in the model.28 The solution pool feature of CPLEX is used to generate a set of distinct starting points (250 for the root node, 10 for all others) for the original model. The binary variables are fixed at each starting point, and the resulting nonlinear optimization model (NLP) is solved using CONOPT.88 The solution of the model is updated only if the solution of the NLP is lower than the existing upper bound. The nodes with a lower bound within ε tolerance from the best upper bound (LBnode/UB ≥ 1 − ε) are eliminated from the branch-and-bound tree. A more detailed explanation of branch-and-bound algorithms are available in the textbooks of Floudas89,90 and reviews of the global optimization methods.91−93

engineering, contingency, startup, royalties, fees, and spare parts.51 The ICs are assumed to be 32% of the TDC. In order to effectively compare various CBTL technologies, it is imperative to levelize production costs. The methodology proposed by Kreutz et al.51 is utilized to calculate the capital charges (CC) by multiplying the levelized capital charge rate (LCCR) and the interest during construction factor (IDCF) by the total overnight capital as shown in eq 4: (4)

CC = LCCR· IDCF· TPC

Kreutz calculates a LCCR value of 14.38%/y and an IDCF value of 7.16%/y. An overall multiplier of 15.41%/y is used to convert the TPC into total capital charges. Additionally, the plant is assumed to operate with a capacity (CAP) of 330 days/y and operating/maintenance (OM) costs are assumed to be 4.5% of the TPC. Using these assumptions, the total levelized cost of a unit (CostU) is calculated using eq 5: costUu =

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

3. COMPUTATIONAL STUDIES The process synthesis model described in the previous sections (see also Supporting Information where the complete model is presented) was used to examine two sets of studies utilizing three different combinations of feedstocks. A global optimization framework was utilized in each case study and terminated if all nodes in the branch-and-bound tree were processed or if 100 CPU hours had passed.28 The first set included a total of 12 case studies that input hardwood biomass and LV bituminous coal. In order to show the effect of the refinery capacity, four scales of 1, 5, 10, and 50 thousand barrels per day (kBD) of gasoline equivalent (based on the lower heating value of gasoline) were investigated. Three sets of liquid fuels products were considered that represented ratios (a) commensurate with the 2012 United States demand (i.e., 67 vol % gasoline, 22 vol % diesel, 11 vol % kerosene),94 (b) that maximized diesel production, or (c) that freely output the liquid products. The case studies are denoted as N − C, where N represents the fuel composition of gasoline/ diesel/kerosene in the unrestricted (U), maximize diesel (D), or U.S. ratios (R) cases and C represents the refinery capacity. For example, U-5 represents a refinery that outputs 5 kBD of gasoline equivalent in any unrestricted composition. A 70% minimum carbon input threshold from coal was imposed for the 12 case studies utilizing hardwood and LV bituminous coal. Using the GHG emissions of a petroleum-based refinery (91.6 kg CO2eq/GJ LHV)53 and those of a natural gas combined cycle plant (101.3 kg CO2eq/GJ) that produces electricity,68 at least a 50% reduction in the GHG emissions from these processes are imposed in the first set of case studies. Thus, the global optimization algorithm guarantees that the optimal solution has life-cycle GHG emissions that are at most 50% of these fossil fuel-based processes. Note that carbon capture and sequestration are not utilized in the calculation of the GHG emissions for the natural gas combined cycle.68 The GHG emissions from electricity are added to the life-cycle greenhouse gas emissions of the refinery if electricity is input and subtracted from the life-cycle greenhouse gas emissions if electricity is output from the CBTL plant. The second set includes six case studies that input LV bituminous coal with duckweed biomass and six case studies that cofeed coal from Anhui, China, with duckweed biomass. Three scales of 1.6, 5, and 10 thousand barrels per day (kBD) of gasoline equivalent are investigated and two sets of liquid fuels products that maximize the production of diesel (D) and freely

(5)

where the production (Prod) is used to levelize the cost based on the total energy of products produced. 2.9. Objective Function. The objective function of the MINLP model that is minimized is given by eq 6: F F El MINcostFBio + cost Coal + cost OF H 2 + costBut + cost

+ costSeq − sales LPG +

∑ u ∈ UInv

costUu (6)

where the feedstock costs are as follows: costFBio is the biomass cost, costFCoal is the coal cost, costHF 2O is the input water cost, costFBut is the input butane cost. The electricity cost is denoted by costEl and is negative if the CBTL refinery produces electricity as a byproduct. CostSeq is the CO2 sequestration cost, costU is the levelized investment cost, and salesLPG is the profit from the sale of LPG. Each term in the objective function is normalized with respect to the total energy of products. The authors note that other normalization factors (total volume of products, etc.) and other objective functions (maximizing net present value, etc.) can also be incorporated into the process synthesis framework. The process synthesis model, along with simultaneous heat, power, and water integration, that was summarized in the previous sections forms a large-scale nonconvex mixed-integer nonlinear optimization (MINLP) model which is solved to global optimality using a branch-and-bound algorithm.28 The overall model has 15,545 continuous variables, 37 binary variables, 19,418 constraints, and 390 nonconvex terms, which consist of 327 bilinear terms, 3 quadrilinear terms, and 60 power functions. At every node in the tree, the nonlinear equations are replaced with their linear relaxations, and a mixed-integer linear problem (MILP) is solved via CPLEX.87 That node is then branched into two children nodes. The linear relaxation of the nonlinear equations are formed by a piecewise-linear underestimating scheme which introduces binary variables, depending logarithmically on the number of pieces to be partitioned. The existence of nonconvex bilinear terms is related to flash units whose vapor−liquid phase equilibrium must be specified, streams with unknown compositions that must be split, or when units within the CBTL superstructure require chemical equilibrium to be enforced. The number of partitions selected for the flash units was equal to 8. The number of partitions selected for units that are restricted by the water-gas-shift L



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output any composition of liquid products (U) are considered. Each case study will input 3 times as much coal as biomass (due to biomass availability considerations and within a 3% confidence interval) on an as-received basis and at least a 30% reduction in GHG emissions is imposed. Table 3 lists the cost parameters used in the process synthesis model. The costs related to the delivery of the feedstock are included in the feedstock (biomass, coal, butanes, and freshwater) costs, whereas the product costs (i.e., electricity and propane) do not include any transportation costs to the consumer. The costs associated with the capture and compression of CO2 are included in the investment costs of the plant and the cost for the transportation, storage, and monitoring of CO2 is displayed in Table 3.29

The optimal process synthesis topology will include the minimum number of heat exchanger matches based on (i) the fluid flow rates and operating conditions of the heat engine, (ii) the amount of electricity produced by the heat engines, (iii) the amount of cooling water required by the heat engines, and (iv) the location of all pinch points in the subnetworks.26,27,89,95 As described in previous works by Floudas et al. and Elia et al. the minimum annualized cost can then be calculated.25,89,95 The investment cost of the heat exchanger network is then added to the overall CBTL refinery investment cost.29 3.1. Optimal Process Topologies. 3.1.1. LV Bituminous Coal and Hardwood Biomass. The optimal process topologies selected within the CBTL refinery for the 12 case studies utilizing hardwood and LV bituminous coal are shown in Table 4. Biomass and coal are converted into synthesis gas using either solid or a combination of solid and recycle gases input into the gasifier. The biomass gasifier operated using a solid/vapor fuel for all 12 case studies, while the coal solid/vapor fueled gasifier was selected for all case studies except U-1. The benefit of utilizing a combination of solid/vapor fuel into the gasifiers is based on its ability to convert some of the CO2 in the CBTL refinery into additional synthesis gas. Three temperatures were considered for operation of the biomass gasifier (900 °C, 1000 °C, or 1100 °C) and coal gasifier (1100 °C, 1200 °C, or 1300 °C). In all 12 case studies, the

Table 3. Cost Parameters (2012 $) for the CBTL Refinery29 item

cost

hardwood

$70.00/dry metric ton $50.00/dry metric ton $1.84/gal $0.07/kW h $0.50/metric ton

duckweed butanes electricity freshwater a

item

cost

LV bituminous Anhui, China, coal propanes CO2 TS&Ma

$93.41/dry short ton $56.78/dry short ton $1.78/gal $5/metric ton

TS&M: transportation, storage, and monitoring.

Table 4. Topological Information for the Optimal Solutions for the 24 Case Studies Is Showna LV bituminous and hardwood biomass case study

U-1

U-5

D-1

D-5

D-10

D-50

R-10

R-50

biomass conv. biomass temp. coal conv. coal temp WGS/RGS temp. min wax FT nom. wax FT FT upgrading MTG usage MTOD usage CO2SEQ usage GT usage

S/V 900 S 1100 300 − − − Y − Y −

S/V S/V S/V S/V 900 900 900 900 S/V S/V S/V S/V 1100 1100 1100 1100 300 300 300 300 − − − − − − − − − − − − Y Y Y − − − − Y Y Y Y Y − − − − LV bituminous and duckweed biomass

U-10

U-50

S/V 900 S/V 1100 300 − − − − Y Y −

S/V 900 S/V 1100 300 − − − − Y Y −

S/V S/V S/V S/V 900 900 900 900 S/V S/V S/V S/V 1100 1100 1100 1100 300 300 300 300 − − − − − − − − − − − − − − Y Y Y Y Y Y Y Y Y Y − − − − Anhui, China, coal and duckweed biomass

R-1

R-5

S/V 900 S/V 1100 300 − − − Y Y Y −

case study

U-1.6

U-5

U-10

D-1.6

D-5

D-10

U-1.6

U-5

U-10

D-1.6

D-5

D-10

biomass conv. biomass temp. coal conv. coal temp WGS/RGS temp. min wax FT nom. wax FT FT upgrading MTG usage MTOD usage CO2SEQ usage GT usage

S 1000 S/V 1100 300 − − − Y − Y −

S/V 900 S/V 1300 300 − − − Y − Y −

S/V 1100 S/V 1300 300 − − − Y − Y −

S 900 S/V 1100 300 − − − − Y Y −

S 1100 S/V 1100 300 − − − − Y Y −

S 1100 S/V 1100 300 − − − − Y Y −

S 900 S/V 1100 300 − − − Y − Y −

S 1100 S/V 1100 300 − − − Y − Y −

S 1100 S/V 1200 300 − − − Y − Y −

S 900 S/V 1100 300 − − − − Y Y −

S 1100 S/V 1200 300 − − − − Y Y −

S 900 S/V 1200 300 − − − − Y Y −

a

Biomass and coal conversion (biomass conv., coal conv., respectively) is gasification with a solid (S) or solid/vapor (S/V) fueled system. Temperatures (temp.; °C) of the biomass and coal gasifiers, along with the operating temperature of the forward or reverse water-gas-shift reactor, are shown. The presence of a gas turbine (GT) or a CO2 sequestration system (CO2SEQ) is noted using yes (Y) or no (−). The FT units will be designated as either cobalt-based or iron-based. The FT hydrocarbons will be upgraded using fractionation (fract.) or ZSM-5 catalytic conversion. Methanol conversion using methanol-to-gasoline (MTG) and methanol-to-olefins/olefins-to-gasoline-and-diesel (MTO/MOGD) is noted using yes (Y) or no (−). M



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constant. In all the case studies, the operating temperature of the dedicated WGS reactor was selected to be 300 °C. The three unrestricted case studies selected syngas conversion to methanol and then the methanol-to-gasoline technology as the optimal route for liquid fuels production. The diesel case studies also utilized methanol conversion, followed by the methanol-to-olefins and olefins-to-gasoline/distillate technology to output a combination of gasoline, diesel, and kerosene. CO2 sequestration was necessary in order to meet the 30% reduction in GHG emissions imposed for this set of case studies. 3.1.3. Anhui, China, Coal and Duckweed Biomass. Table 4 also illustrates the optimal process topologies for the six case studies that input coal from Anhui, China, and duckweed biomass. None of the case studies utilize a recycle gas feed to the biomass gasifier and instead only input solid duckweed. On the contrary, all of the coal gasifiers operated using a combination of vapor and solid feed. The operating temperature of the biomass gasifiers was 1100 °C for the U-5, U-10, and D-5 cases and 900 °C for the remaining case studies (U-1.6, D-1.6, and D-10). The U-1.6, U-5, and D-1.6 case studies utilized a 1100 °C coal gasifier, while the U-10, D-5, and D-10 case studies used a 1200 °C coal gasifier. A dedicated water-gas-shift reactor was selected in each of the case studies and operated at 300 °C. The unrestricted case studies all utilized a methanol intermediate that selected the MTG technology to produce gasoline. The maximum diesel case studies also utilized methanol conversion that was followed by the MTO/MOGD processes. These topological selections are consistent with the ones selected in the previous case studies. CO2 sequestration was required in each of these case studies to meet the 30% reduction in GHG emissions imposed for this set of case studies. 3.2. Overall Costs of Liquid Fuels. Table 5 illustrates the overall cost results for the 24 case studies. The total cost is the sum of the feedstock costs, the CO 2 sequestration costs, the investment costs, and the operating and maintenance costs. Electricity and LPG may be sold as byproducts and are subtracted from the total cost of the CBTL refinery. The figures in Table 5 are normalized with respect to the total energy (GJ) of products. The break even oil price (BEOP) is calculated using the refiner’s margin26,51 and is illustrated in dollars per barrel ($/bbl); it represents the price at which the CBTL refinery becomes competitive with petroleum-based processes. The lower bound, in $/GJ, is shown for the optimal solution. The corresponding optimality gap is calculated for each of the case studies and is less than 10% in all cases. Additionally, the number of nodes that were evaluated in the branchand-bound tree and the time it took for the branch-and-bound step in the global optimization framework are shown in Table 5. 3.2.1. LV Bituminous Coal and Hardwood Biomass. As shown in Table 5, costs associated with the capital investment (capital charges and O&M) are the largest contributor to the overall cost of the CBTL refinery for the LV bituminous and hardwood case studies. A singular train is needed for most sections of the CBTL plant when the capacity is less than 10 kBD; therefore, a significant economy of scale is expected. Since the refinery will only require one or two units to produce the desired amount of fuels at low scales, the capital charges diminish appreciably when comparing the 1 kBD and 10 kBD scales, and thus, the refinery can take full advantage of the low scaling factor (0.5−0.7). However, as the refinery scale grows beyond 10 kBD, multiple units will need to be introduced in order to produce the desired amount of fuels. Thus, the scaling factor for these units will begin to approach 0.9,51 and the benefits from the economies of scale will begin to dwindle. The BEOP ranges from

operating temperature of the biomass gasifier was selected to be 900 °C, while the coal gasifier operated at 1100 °C for all case studies. Gasifiers operating at lower temperatures have less favorable conditions for CO2 consumption because of the higher value of the forward water-gas-shift equilibrium rate constant. In addition, these gasifiers will produce less waste heat for steam generation. However, investment and utility costs for oxygen generation may be decreased since these units will require less O2 for combustion. Higher temperatures have more favorable conditions for CO2 consumption and generate more waste heat for steam generation. Four temperatures were considered for the operation of the dedicated water-gas-shift reactor (300 °C, 400 °C, 500 °C, or 600 °C). In all case studies, the operating temperature of the dedicated WGS unit was chosen to be 300 °C in order to increase the H2/CO ratio of the synthesis gas for hydrocarbon production at the expense of producing more CO2. The type of liquid fuels produced by the CBTL refinery dictated the selection of the hydrocarbon production and upgrading route chosen. The four unrestricted case studies all utilized methanol synthesis and methanol-to-gasoline (MTG) as the optimal technology. This is due to the lower capital cost associated with the MTG route over the FT routes. The four case studies that maximized diesel output selected the methanol-toolefins (MTO) and olefins-to-gasoline/distillate (MOGD) technologies. The case studies that produced fuels commensurate with U.S. demand displayed a topological switch between the smallest capacities. At a capacity of 1 kBD, the MTO/MOGD process was the optimal topology, since the Mobil olefins-togasoline/distillate technology can produce various ratios of gasoline, diesel, and kerosene depending on which mode it operates in. As the refinery scale increased, the optimal topology shifted to include MTG along with MTO/MOGD. The MTG unit will produce the majority of the gasoline, while the MTO/ MOGD unit will produce the balance of diesel and kerosene. CO2 sequestration was selected in the 12 case studies in order to meet the 50% reduction requirement from petroleum based processes. The methane- and ethane-rich gases generated from the CBTL refinery were also split between the fuel combustor and the autothermal reformer. The reformed gases from the autothermal reactor are recycled back to the gasifiers. The fuel combustor provides the necessary heat required for the biomass drier to reduce the moisture content of hardwood from 45 to 20 wt %. None of the case studies utilized a gas turbine. 3.1.2. LV Bituminous Coal and Duckweed Biomass. Table 4 shows the optimal process topologies for the six case studies utilizing LV bituminous coal and duckweed biomass in the CBTL refinery. The biomass gasifier operates using a solid fuel for all case studies except for U-5 and U-10, where it also inputs recycle gases. The biomass gasifier operates at 1100 °C in the U-10, D-5, and D-10 case studies, 1000 °C in the U-1.6 case study, and 900 °C in the U-5 and D-1.6 case studies. The coal solid/vapor fueled gasifier was selected for all six of the case studies. For all case studies except for U-5 and U-10, the coal gasifier operates at 1100 °C. In the U-5 and U-10 case studies, the coal gasifier operated at 1300 °C. Gasifiers that operate at lower temperatures require less oxygen which reduces the steam and electricity requirements from the CBTL refinery. Additionally, these units will produce less waste heat from syngas cooling and may have larger levels of CO2 in the effluent since the equilibrium constant of the forward water-gas-shift reaction is higher. Gasifiers operating at higher temperatures will require more oxygen; however, they may reduce the amount of CO2 in the effluent due to the lower value of the forward water-gas-shift equilibrium N



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Table 5. Overall Cost Results for the 24 Case Studiesa LV bituminous and hardwood biomass contribution to cost ($/GJ of products)

U-1

biomass coal butane water CO2 TS&M investment O&M electricity LPG total ($/GJ) total ($/bbl) lower bound ($/GJ) gap (%) nodes branch and bound time (s)

1.86 3.96 0.00 0.02 0.37 17.15 4.53 1.68 −2.53 27.03 136.01 26.71 1.19 49 14 300

U-5

U-10

U-50

D-1

1.97 2.03 2.03 1.60 4.19 4.33 4.33 3.40 0.00 0.00 0.00 0.00 0.02 0.02 0.02 0.02 0.40 0.41 0.41 0.33 10.10 8.13 6.17 15.96 2.67 2.15 1.63 4.21 0.75 0.26 0.26 2.20 −2.53 −2.53 −2.53 0.00 17.56 14.79 12.32 27.72 83.89 68.65 55.03 139.82 16.81 14.09 11.82 26.89 4.30 4.76 4.07 3.02 167 300 179 15 230 000 295 000 362 000 29 000 LV bituminous and duckweed biomass

D-5

D-10

1.78 3.78 0.00 0.02 0.37 9.67 2.55 0.61 0.00 18.79 90.63 17.83 5.10 80 209 000

1.83 3.91 0.00 0.02 0.39 7.78 2.05 0.15 0.00 16.13 76.13 15.26 5.44 189 361 000

D-50

R-1

R-5

R-10

1.84 1.58 1.89 1.95 3.91 3.37 4.03 4.16 0.00 0.00 0.00 0.00 0.02 0.02 0.02 0.02 0.39 0.33 0.39 0.40 5.63 15.85 10.29 8.27 1.49 4.19 2.72 2.18 0.15 2.22 0.70 0.22 0.00 0.00 −1.53 −1.53 13.42 27.56 18.49 15.67 61.11 138.95 88.98 73.54 12.10 26.87 17.47 14.79 9.87 2.50 5.52 5.63 218 221 300 300 361 000 271 000 182 000 219 000 Anhui, China, coal and duckweed biomass

R-50 1.95 4.16 0.00 0.02 0.40 6.03 1.59 0.22 −1.53 12.85 57.94 12.29 4.38 300 202 000

contribution to cost ($/GJ of products)

U-1.6

U-5

U-10

D-1.6

D-5

D-10

U-1.6

U-5

U-10

D-1.6

D-5

D-10

biomass coal butane water CO2 Seq investment O&M electricity LPG total ($/GJ) total ($/bbl) lower bound ($/GJ) gap (%) nodes branch and bound time (s)

0.76 4.91 0.00 0.02 0.43 15.61 4.12 1.18 −2.53 24.49 122.03 22.45 8.34 522 360 000

0.78 5.08 0.00 0.03 0.44 10.65 2.81 0.59 −2.53 17.85 85.48 16.13 9.65 300 290 000

0.79 5.11 0.00 0.03 0.45 8.40 2.22 0.52 −2.53 14.99 69.72 13.53 9.71 120 65 900

0.67 4.35 0.00 0.02 0.40 14.73 3.89 1.39 0.00 25.45 127.31 23.84 6.33 300 124 000

0.70 4.57 0.00 0.02 0.43 10.08 2.66 0.76 0.00 19.22 93.00 17.74 7.69 300 106 000

0.72 4.65 0.00 0.02 0.43 7.98 2.11 0.55 0.00 16.45 77.78 15.42 6.28 300 201 000

0.92 3.64 0.00 0.02 0.37 15.98 4.22 0.91 −2.53 23.53 116.73 22.45 4.58 522 360 000

0.95 3.76 0.00 0.02 0.38 10.91 2.88 0.45 −2.53 16.83 79.85 15.81 6.09 475 360 000

0.97 3.81 0.00 0.02 0.39 8.76 2.31 0.28 −2.53 14.01 64.32 13.02 7.05 700 336 000

0.84 3.30 0.00 0.02 0.36 15.40 4.07 0.84 0.00 24.82 123.86 23.06 7.10 388 73 400

0.86 3.40 0.00 0.02 0.37 10.54 2.78 0.32 0.00 18.28 87.85 17.08 6.58 700 272 000

0.87 3.42 0.00 0.02 0.37 8.31 2.20 0.36 0.00 15.55 72.79 14.46 6.97 700 164 000

a

The contributions to the total costs (in $/GJ) come from biomass, coal, butane, water, CO2 transportation/storage/monitoring (CO2 TS&M), investment, input electricity, and operations/maintenance (O&M). LPG and output electricity are allowable byproducts (negative value). The overall costs, as well as the lower bound values, are reported and the optimality gap between the optimal solution and the lower bound is shown. The number of nodes processed in the branch-and-bound algorithm, as well as the solution times for the branch-and-bound step, are illustrated.

$136/bbl−$140/bbl for a 1 kBD plant, $83/bbl−$91/bbl for a 5 kBD plant, $68/bbl−$77/GJ/bbl for a 10 kBD plant, and $55/bbl−$62/bbl for a 50 kBD plant. The overall cost of the CBTL refinery depends on the composition of liquid fuels produced. The lowest overall cost is achieved when no restriction is placed on the liquid products; the optimal technology selected is methanol synthesis and, subsequently, methanol-to-gasoline. The finished liquid fuels composition from the MTG reactor consists of 82 wt % gasoline and 10 wt % LPG. The production and profit from selling byproduct LPG provides additional economic incentive for the refinery. The United States ratio case studies provide the next lowest cost of fuels production. At scales larger than or equal to 5 kBD, the optimal topology includes both MTG and MTO/ MOGD. LPG is produced as a consequence of the selection of the MTG reactor and is sold as a byproduct. Finally, the maximum diesel case studies provide the largest cost of fuel production. This is a direct consequence of the type of technology chosen to produce the desired composition of products. The lower cost associated with the R-1 case study, when compared to

D-1, is due to the minor difference in the amount of feedstock necessary to produce the desired amount of liquid fuels. The slightly smaller flow rate in the R-1 case study ultimately leads to lower capital charges since smaller units can meet the production constraint. 3.2.2. LV Bituminous Coal and Duckweed Biomass. Table 5 also shows the overall cost results for the six case studies utilizing LV bituminous coal and duckweed biomass. The capital investment contributes the largest portion to the total cost of the CBTL refinery. The levelized unit investment cost represents 64% of the total cost CBTL refinery cost in the U-1.6 case and 56% of the total cost in the U-10 case. In the maximum diesel case studies, the levelized unit investment cost represents 58%, 52%, and 49% of the total CBTL cost for the D-1.6, D-5, and D-10 refineries, respectively. The unrestricted case studies had a lower overall cost than the maximum diesel case studies, due to the byproduct LPG that offset the total cost of the refinery. The BEOP ranges from $122/bbl−$127/bbl for a 1.6 kBD plant, $85/bbl−$93/bbl for a 5 kBD plant, and $69/bbl−$78/GJ/bbl for a 10 kBD plant. 3.2.3. Anhui, China, Coal and Duckweed Biomass. The overall cost results for the Anhui, China, coal and duckweed O



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increases, the feedstock costs represent a larger portion of the overall cost of liquid fuels production due to economies of scale. In order to examine the effect of the feedstock price on the BEOP, the BEOP is calculated for the U-10 case studies, assuming the biomass and coal are priced ±$10.00/dry ton from their nominal values. Figure 13 illustrates the results from the parametric analysis. In Figure 13a, the price of coal is varied while the biomass costs stay constant. Figure 13b shows the effect of varying the biomass price on a constant coal price. The break even oil price (BEOP) ranges from $60/barrel to $73/barrel, depending on the prices of coal and biomass, types of coal and biomass, and GHG reduction constraints. The Anhui, China, coal and duckweed combination, which has a 30% reduction in GHG emissions from petroleumbased processes imposed, has the lowest BEOP. The U-10 refinery that inputs LV bituminous coal and duckweed biomass, and has a 50% reduction in GHG emissions imposed, has the largest BEOP. Varying the coal price has a more pronounced effect on the BEOP because of the higher fraction of coal input into the refinery relative to biomass. This is evident from the steeper slopes in Figure 13a. 3.4. Investment Costs. 3.4.1. LV Bituminous Coal and Hardwood Biomass. Table 6 shows the total plant cost (TPC) of the CBTL refinery, as well as the breakdown for the major sections of the plant, for the LV bituminous coal and hardwood biomass case studies. The major sections include syngas generation, syngas cleaning, hydrocarbon production, hydrocarbon upgrading, hydrogen/oxygen production, heat and power integration, and wastewater treatment. In each of the case studies, the syngas generation section contributes the largest portion toward the total plant cost of the CBTL refinery. The syngas generation section includes the costs associated with the handling of the feedstocks as well as the costs of the coal and biomass gasifiers. The second largest contributor to the total plant cost is the syngas cleaning section. The cost of the dedicated watergas-shift reactor, as well as CO2 capture and compression, are included in this cost. Syngas generation represents between 34 and 40% of the total plant cost, while the syngas cleaning represents about 21−24% of the TPC. The third largest contributor to the TPC is the hydrocarbon upgrading section, which consists of the methanol conversion units as well as all units associated with the FT hydrocarbon upgrading route. This section represents between 11 and 17% of the TPC. The hydrocarbon production section and the hydrogen/oxygen production section contribute roughly the same amount toward the total plant cost. The hydrocarbon production section consists of the FT units as well as the methanol synthesis reactor and represents about 7−11% of the TPC. The hydrogen/oxygen production section represents 8−10% of the TPC. The heat and power integration section consists of the heat exchangers in the refinery and contributes 4−9% to the overall cost. Finally, the wastewater treatment section represents between 2 and 4% of the TPC. The total plant cost can be expressed as an “overnight” cost by factoring in the preproduction costs, inventory capital, financing costs, and other owners’ costs.51,68 As Table 6 illustrates, the TPC ranges from $187−$202 MM for 1 kBD plants, $570−$607 MM for 5 kBD plants, $917−$975 MM for 10 kBD plants, and $3318−$3636 MM for 50 kBD plants. The normalized cost of the production is shown by dividing the TPC by the production capacity of the CBTL refinery. The economy of scale associated with the CBTL refinery is illustrated when comparing the normalized values. For example, when scaling the refinery from 1 kBD to 5 kBD, there is between a 35 to 41% decrease in the normalized capital cost.

biomass are also shown in Table 5. The overall cost of these CBTL refineries is lower than the ones utilizing LV bituminous coal and duckweed biomass. This is because the cost of coal from Anhui, China ($56.78/dry short ton) is less expensive than LV bituminous coal ($93.41/dry short ton). The investment costs associated with the Anhui, China, coal case studies are slightly higher due to the larger ash content of this coal relative to that of LV bituminous coal, which requires larger units to meet the production constraints. The cost of feedstocks represents between 19 and 34% of the total refinery cost for the unrestricted case studies and 16−28% of the total refinery cost for the maximum diesel case studies. The levelized unit investment cost represents 68%, 65%, and 62% of the total CBTL cost for the U-1.6, U-5, and U-10 case studies, respectively. The levelized unit investment costs for the case studies that maximized diesel represent 62%, 58%, and 53% of the total cost for capacities of 1.6 kBD, 5 kBD, and 10 kBD, respectively. The BEOP ranges from $116/bbl−$124/bbl for a 1.6 kBD plant, $79/bbl−$88/bbl for a 5 kBD plant, and $64/bbl−$73/GJ/bbl for a 10 kBD plant. 3.2.4. Cost Comparison. This section will compare the overall costs for the three feedstock combinations investigated. The comparison will be divided between the case studies which output an unrestricted composition and those that maximize the diesel product. Unrestricted Case Studies. Table 5 shows that the Anhui, China, coal and duckweed biomass combination is the most profitable for outputting liquid fuels in any composition. For the U-5 and U-10 case studies, the break even oil price is $80 and $64, respectively. For similar capacities, the break even oil price for systems utilizing LV bituminous coal and duckweed biomass is $85 and $70, respectively. The Anhui, China, coal and duckweed biomass combinations input more coal and biomass into the refinery than the systems utilizing LV bituminous and duckweed biomass, thus leading to higher investment costs. However, the cost of coal from Anhui, China, is significantly less expensive than LV bituminous, which leads to the differences in the overall costs between these two systems. These case studies achieve a 30% reduction in life-cycle greenhouse gas emissions compared to fossil-fuel based processes. Systems utilizing LV bituminous coal and hardwood have a BEOP of $84 for a 5 kBD plant and $69 for a 10 kBD plant. However, these systems achieve a 50% reduction in life-cycle emissions, as illustrated later in the text. Maximization of Diesel Case Studies. The Anhui, China, coal and duckweed biomass combination is also the most profitable feedstock combination for maximizing the production of diesel. For the D-5 and D-10 case studies, the break even oil price is $88 and $73, respectively. For similar capacities, the break even oil price for systems utilizing LV bituminous coal and duckweed biomass is $93 and $78, respectively. The Anhui, China, coal and duckweed refineries have larger investment costs. However, the contribution from coal is significantly cheaper for systems utilizing Anhui, China, coal, and this difference makes it more attractive than utilizing LV bituminous. For plants utilizing LV bituminous coal and hardwood, the BEOP is $91 for a 5 kBD plant and $76 for a 10 kBD plant. 3.3. Parametric Analysis. Table 5 displays the BEOP based on a coal-delivered price of $93.41/dry short ton for LV bituminous coal and $56.78/dry short ton for Anhui, China, coal and a biomass delivered price of $70.00/dry ton for hardwood and $50.00/dry ton for duckweed. In the 24 case studies examined, the feedstock costs represent between 16 and 52% of the total costof the CBTL refinery. In addition, as the refinery size P



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Figure 13. Parametric analysis of the coal and biomass cost. The break even oil price (BEOP) is plotted for the case studies with an unrestricted product composition and a 10 kBD capacity (U-10). The effect of coal price (with nominal biomass prices) is shown in (a), while the effect of biomass price (with nominal coal prices) is shown in (b).

3.4.2. LV Bituminous Coal and Duckweed Biomass. The TPC for the case studies utilizing LV bituminous coal and duckweed biomass are also shown in Table 6. Syngas generation, the largest contributor to the TPC, represents between 33 and 35% of the total plant cost. Syngas cleaning, hydrocarbon production, and hydrocarbon upgrading represent between 22−23%, 8−9%, and 12−13% of the total plant cost, respectively. Hydrogen and oxygen production represents between 8 and 10% of the total plant cost, while heat and power integration contributes about 8−11% to the TPC. Wastewater treatment, the smallest contributor to the TPC, contributes about 3% to the total plant cost. For a 1.6 kBD plant, the TPC ranges between $278−$294 MM. The TPC for a 5 kBD plant is between $594−$628 MM, while a 10 kBD plant has a TPC between $940−$991 MM. 3.4.3. Anhui, China, Coal and Duckweed Biomass. Table 6 shows the total plant cost (TPC) for the CBTL refineries inputting coal from Anhui, China, and duckweed biomass. Syngas generation, syngas cleaning, hydrocarbon production, and hydrocarbon upgrading represent between 36−39%, 21− 22%, 8−9%, and 12−13% of the TPC, respectively. Additionally, hydrogen and oxygen production, heat and power integration, and wastewater treatment represent between 7−8%, 7−9%, and 3% of the total plant cost, respectively. The total plant cost of these refineries is larger than those utilizing LV bituminous coal and duckweed biomass because of the increased cost for

syngas generation. The increased cost of the syngas generation section is due to the larger flow rate of coal through the units. As shown in Table 1, coal from Anhui, China, has about 4 times the ash content as that of LV bituminous. The larger ash content of coal from Anhui, China, means that the CBTL refinery needs to input more coal to meet the production requirement of the plant. The TPC ranges from $290 to $301 MM for 1.6 kBD plants, $621 to $643 MM for 5 kBD plants, and $980 to $1033 MM for 10 kBD plants. 3.5. Material and Energy Balances. 3.5.1. LV Bituminous Coal and Hardwood Biomass. Material balances for the 12 cases studies utilizing LV bituminous coal and hardwood biomass are shown in Table 7. Biomass and coal are shown in dry tons per hour (dt/h), while butane, water, and the liquid products are shown in thousand barrels per day (kBD). Sequestered and vented CO2 are shown in tons per hour. For the unrestricted case studies, gasoline is the only liquid fuel produced through the MTG process. The refinery also produces about 19% vol LPG. The U.S. ratios case studies produce gasoline, diesel, and kerosene. In the R-1 case, the MTO and MOGD processes were chosen as the optimal process technology, and therefore no LPG is produced. For the larger U.S. ratios case studies, MTG is incorporated in the optimal process topology, so that the refinery produces an additional 11% vol of LPG. The maximum diesel case studies also produced all three liquid fuel products. Diesel Q



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Table 6. Breakdown of the Investment Costs for the 24 Case Studiesa LV bituminous and hardwood biomass contribution to cost (MM $) syngas generation syngas cleaning hydrocarbon production hydrocarbon upgrading hydrogen/oxygen production heat and power integration wastewater treatment total (MM $) total ($/bpd) contribution to cost (MM $) syngas generation syngas cleaning hydrocarbon production hydrocarbon upgrading hydrogen/oxygen production heat and power integration wastewater treatment total (MM $) total ($/bpd)

U-1

U-5

U-10

217 136 48 79 50

11 7 202 202 128

44 80 240 8 20 33 93 7 595 958 3636 188 119 049 95 817 72 720 206 798 LV bituminous and duckweed biomass U-5

U-10

1404 800 297 513 289

D-1

71 48 18 28 20

U-1.6

354 216 74 123 78

U-50

D-1.6

66 45 17 27 19

D-5

D-5

D-10

D-50

R-1

334 204 71 121 76

1299 741 276 397 281

44 20 570 125 338

79 32 917 100 748

235 7 44 80 91 7 20 32 3318 187 607 975 72 884 194 744 125 693 100 895 Anhui, China, coal and duckweed biomass

U-1.6

U-5

U-10

213 133 47 99 51

R-10

205 129 46 77 50

D-10

65 45 17 27 19

R-5

D-1.6

346 211 73 155 78

D-5

R-50 1361 776 288 514 288 238 92 3556 73 597 D-10

99 69 25 38 29

218 146 52 79 58

350 228 82 124 88

91 64 24 37 27

200 135 47 77 52

320 211 72 121 78

112 68 27 38 26

247 142 54 79 50

407 223 86 124 76

105 64 25 37 25

233 135 53 77 51

372 209 79 121 71

25 10 294 183 982

53 22 628 125 534

85 34 991 99 081

26 9 278 191 577

65 19 594 130 602

107 30 940 103 322

21 10 301 188 398

50 21 643 128 666

83 34 1033 103 293

26 9 290 200 342

51 21 621 136 521

96 32 980 107 692

a

Costs of the syngas generation, syngas cleaning, hydrocarbon production, hydrocarbon upgrading, hydrogen/oxygen production, heat and power integration, wastewater treatment sections of the CBTL refinery are shown. All values are shown in MM$ and are also expressed in $/bpd by dividing the total investment cost by the total plant capacity.

Table 7. Overall Material Balance for the 24 Case Studiesa LV bituminous and hardwood biomass material balances

U-1

U-5

U-10

U-50

D-1

biomass (dt/h) coal (dt/h) butane (kBD) water (kBD) gasoline (kBD) diesel (kBD) kerosene (kBD) LPG (kBD) seq CO2 (tonne/h) vented CO2 (tonne/h)

6.09 8.28 0.00 1.31 1.00 0.00 0.00 0.19 17.00 3.07

material balances

U-1.6

U-5

U-10

D-1.6

D-5

D-10

U-1.6

U-5

U-10

D-1.6

D-5

D-10

biomass (dt/h) coal (dt/h) butane (kBD) water (kBD) gasoline (kBD) diesel (kBD) kerosene (kBD) LPG (kBD) seq CO2 (tonne/h) vented CO2 (tonne/h)

5.56 17.39 0.00 2.71 1.60 0.00 0.00 0.30 31.45 3.78

17.97 56.21 0.00 9.06 5.00 0.00 0.00 0.93 102.06 14.81

36.17 113.17 0.00 18.36 10.00 0.00 0.00 1.86 206.00 30.44

4.92 15.40 0.00 2.38 0.18 1.09 0.18 0.00 29.49 1.38

16.16 50.54 0.00 7.94 0.56 3.41 0.58 0.00 97.52 7.71

32.89 102.91 0.00 16.34 1.12 6.83 1.15 0.00 199.38 17.70

6.78 21.02 0.00 2.46 1.60 0.00 0.00 0.30 26.84 5.32

21.86 67.74 0.00 8.20 5.00 0.00 0.00 0.93 87.24 19.14

44.31 137.31 0.00 16.71 10.00 0.00 0.00 1.86 177.73 40.30

6.15 19.05 0.00 2.28 0.18 1.09 0.18 0.00 26.09 3.30

19.77 61.26 0.00 7.60 0.56 3.41 0.58 0.00 83.78 13.03

39.76 123.21 0.00 15.26 1.12 6.83 1.15 0.00 169.71 25.89

32.26 66.58 333.29 5.24 43.86 90.54 453.25 7.10 0.00 0.00 0.00 0.00 7.14 15.32 76.92 1.32 5.00 10.00 50.00 0.11 0.00 0.00 0.00 0.68 0.00 0.00 0.00 0.12 0.93 1.86 9.32 0.00 90.73 188.28 943.12 15.21 20.80 47.60 238.70 1.32 LV bituminous and duckweed biomass

D-5

D-10

29.10 39.56 0.00 6.66 0.56 3.41 0.58 0.00 85.38 15.62

60.12 81.76 0.00 14.27 1.12 6.83 1.15 0.00 177.16 36.81

D-50

R-1

R-5

R-10

300.82 5.19 30.96 63.97 409.08 7.05 42.10 86.99 0.00 0.00 0.00 0.00 71.64 1.24 6.96 14.95 5.61 0.64 3.25 6.49 34.15 0.21 1.04 2.08 5.77 0.11 0.54 1.09 0.00 0.00 0.56 1.13 866.77 15.05 88.43 183.67 184.29 1.39 18.88 43.64 Anhui, China, coal and duckweed biomass

R-50 320.18 435.42 0.00 75.03 32.47 10.41 5.44 5.64 919.93 218.67

a

The inputs to the CBTL refinery are coal, biomass, butane, and water, while the outputs include gasoline, diesel, kerosene, LPG, sequestered CO2, and vented CO2.

the 50% reduction of greenhouse gas emissions from petroleumbased processes. Table 8 shows the overall energy balance for each of these case studies. The energy inputs are from the feedstocks and electricity,

was equal to 75% v/v of the total product, kerosene was 13% v/v, and gasoline was 12% v/v of the total liquid fuels product. No byproduct LPG was produced in these case studies. All the case studies sequestered a majority of the CO2 in order to meet R



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Table 8. Overall Energy Balance for the 24 Case Studiesa LV bituminous and hardwood biomass energy balances

U-1

U-5

U-10

U-50

D-1

biomass coal butane gasoline diesel kerosene LPG electricity efficiency (%)

30 78 0 64 0 0 11 6 65.9

160 414 0 319 0 0 57 12 64.0

energy balances

U-1.6

U-5

U-10

D-1.6

D-5

D-10

biomass coal butane gasoline diesel kerosene LPG electricity efficiency (%)

22 164 0 102 0 0 18 6 62.5

70 531 0 319 0 0 57 10 61.5

140 1069 0 637 0 0 113 17 61.2

19 126 0 11 78 13 0 7 59.3

63 478 0 36 243 40 0 12 57.7

128 972 0 71 486 80 0 18 57.0

330 1652 26 855 4282 67 0 0 0 637 3186 7 0 0 49 0 0 8 113 567 0 9 42 7 62.9 62.8 63.5 LV bituminous and duckweed biomass

D-5

D-10

D-50

144 374 0 36 243 40 0 10 60.3

298 772 0 71 486 80 0 5 59.2

1491 3865 0 357 2429 400 0 25 59.2

R-1

R-5

R-10

R-50

26 153 317 67 398 822 0 0 0 41 207 414 15 74 148 8 38 75 0 34 69 7 11 7 63.9 62.7 61.6 Anhui, China, coal and duckweed biomass

1587 4114 0 2069 741 377 343 36 61.5

U-1.6

U-5

U-10

D-1.6

D-5

D-10

26 148 0 102 0 0 18 5 67.2

85 476 0 319 0 0 57 7 66.0

172 965 0 637 0 0 113 9 65.5

24 134 0 11 78 13 0 4 62.9

77 431 0 36 243 40 0 5 62.2

154 866 0 71 486 80 0 12 61.7

a

The energy inputs come from biomass, coal, butane, or electricity; while the energy outputs are gasoline, diesel, kerosene, 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.

67%, while the efficiencies for the maximum diesel case studies are between 61 and 63%. 3.6. Carbon and Greenhouse Gas Balances. 3.6.1. LV Bituminous Coal and Hardwood Biomass. The overall carbon balance (in kg/s) for the CBTL refineries is shown in Table 9. Carbon is input into the refinery by the feedstocks (biomass, coal, or butane) and output as liquid products (gasoline, diesel, or kerosene), LPG byproduct, vented CO2, or sequestered CO2. The amount of carbon input into the refinery by incoming air is negligible and therefore neglected. The trends in Table 9 are consistent with the material balances shown in Table 7. The overall carbon conversion in the 12 case studies utilizing LV bituminous coal and hardwood biomass ranges between 42 and 49%. The amount of carbon leaving as LPG in the unrestricted case studies is constant across the four refinery scales and is about 12% of that leaving as liquid products. The amount of carbon leaving as LPG in the U.S. ratios case studies is around 8% of that leaving as gasoline, diesel, and kerosene. A large portion of the carbon is sequestered in order to meet the 50% reduction in GHG emissions from petroleum-based processes. Table 10 illustrates the total greenhouse gas emissions (in kg CO2 equivalent/s) for the major inputs and outputs of the refinery. A GHG emissions target was enforced for each case study to ensure at most 50% emissions of petroleum-based production of fuels or natural gas-based production of electricity. The total lifecycle GHG emissions (LGHG) is calculated as the sum of the respective emissions from each stage of production. The GHG emission rates include (a) acquisition and transportation of the biomass, coal, and butane feeds, (b) transportation and use of the liquid products and LPG, (c) transportation and sequestration of CO2, (d) venting of CO2, and (e) atmospheric sequestration of CO2 during growth of biomass due to photosynthesis or sequestration. The GHG emissions from (a), (b), (c), (d), and (e) are calculated to compute the well-to-wheel emissions for the refinery from the

while the energy outputs are the liquid products, LPG, and electricity. Each of the values in Table 8 is expressed in MW and is based on the lower heating value of the component. Coal represents the majority of energy input into the refinery in each of the 12 case studies. Electricity is input into the refinery in each of the case studies. The efficiency is calculated by summing the energy contribution of the products and dividing it by the energy contribution of the inputs. Electricity is included in the denominator of the efficiency calculation since it is input into the refinery for each case study. The overall efficiency for each of the 12 case studies ranges from 59 to 66%. 3.5.2. LV Bituminous Coal and Duckweed Biomass. The material and energy balances for the six case studies inputting LV bituminous coal and duckweed biomass are shown in Tables 7 and 8, respectively. The unrestricted case studies only produced gasoline, while the maximum diesel case studies produced gasoline, diesel, and kerosene in a 12:75:13% v/v ratio. A large portion of the CO2 generated in the refinery needed to be captured and sequestered in order to meet the 30% reduction in greenhouse gas emissions. The overall efficiencies of the process are calculated in Table 8 and range between 57 and 59% for the maximum diesel case studies and 61−63% for the unrestricted case studies. 3.5.3. Anhui, China, Coal and Duckweed Biomass. Table 7 also shows the material balances for the case studies utilizing coal from Anhui, China, and duckweed biomass. The flow rates of the feedstocks in these case studies are larger than the flow rates in the LV bituminous and duckweed biomass case studies for corresponding capacities. This is because the higher ash content in the coal from Anhui, China, forces the refinery to input a larger amount in order to meet the required liquid fuels capacity, as explained earlier in the text. The amounts of sequestered and vented CO2 for each case study are also shown. Table 8 illustrates the overall energy balances for these six case studies. The efficiencies for the unrestricted case studies are between 65 and S



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Table 9. Carbon Accounting (in kg/s) for the 24 Case Studiesa LV bituminous and hardwood biomass carbon balances

U-1

U-5

U-10

U-50

biomass coal butane gasoline diesel kerosene LPG vented CO2 seq CO2 % conversion

0.85 1.99 0.00 1.19 0.00 0.00 0.14 0.23 1.29 46.67

carbon balances

U-1.6

U-5

U-10

D-1.6

biomass coal butane gasoline diesel kerosene LPG vented CO2 seq CO2 % conversion

0.60 4.19 0.00 1.90 0.00 0.00 0.22 0.29 2.38 44.33

1.94 13.54 0.00 5.94 0.00 0.00 0.70 1.12 7.74 42.86

3.91 27.26 0.00 11.88 0.00 0.00 1.40 2.31 15.62 42.58

0.53 3.71 0.00 0.21 1.46 0.23 0.00 0.10 2.24 44.92

D-1

D-5

D-10

D-50

4.06 9.53 0.00 0.66 4.58 0.72 0.00 1.18 6.47 43.84

8.38 19.69 0.00 1.32 9.15 1.44 0.00 2.79 13.43 42.43

41.94 0.72 4.32 8.92 98.53 1.70 10.14 20.95 0.00 0.00 0.00 0.00 6.61 0.77 3.85 7.71 45.75 0.28 1.40 2.79 7.19 0.14 0.68 1.35 0.00 0.00 0.42 0.85 13.97 0.11 1.43 3.31 67.23 1.14 6.70 13.92 42.40 48.74 43.91 42.52 Anhui, China, coal and duckweed biomass

D-5

D-10

U-1.6

U-5

U-10

D-1.6

D-5

D-10

1.75 12.17 0.00 0.66 4.58 0.72 0.00 0.58 7.39 42.78

3.56 24.78 0.00 1.32 9.15 1.44 0.00 1.34 15.11 42.03

0.73 3.82 0.00 1.90 0.00 0.00 0.22 0.40 2.03 46.59

2.36 12.32 0.00 5.94 0.00 0.00 0.70 1.45 6.61 45.18

4.79 24.98 0.00 11.88 0.00 0.00 1.40 3.06 13.47 44.57

0.66 3.47 0.00 0.21 1.46 0.23 0.00 0.25 1.98 46.14

2.14 11.15 0.00 0.66 4.58 0.72 0.00 0.99 6.35 44.83

4.30 22.42 0.00 1.32 9.15 1.44 0.00 1.96 12.87 44.59

4.50 9.28 46.47 0.73 10.56 21.81 109.16 1.71 0.00 0.00 0.00 0.00 5.94 11.88 59.38 0.13 0.00 0.00 0.00 0.92 0.00 0.00 0.00 0.14 0.70 1.40 6.98 0.00 1.58 3.61 18.10 0.10 6.88 14.27 71.50 1.15 44.06 42.69 42.64 48.79 LV bituminous and duckweed biomass

R-1

R-5

R-10

R-50 44.64 104.87 0.00 38.54 13.96 6.77 4.23 16.58 69.74 42.47

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

Table 10. Greenhouse Gas (GHG) Balances for the 24 Case Studiesa LV bituminous and hardwood biomass GHG balances

U-1

biomass coal butane gasoline diesel kerosene LPG vented CO2 seq CO2 LGHG GHGAF GHGAE GHGI

−2.76 0.26 0.00 4.28 0.00 0.00 0.52 0.85 0.01 3.16 6.87 −0.56 0.50

U-5

U-10

U-50

D-1

−14.61 −30.16 −150.97 −2.37 1.35 2.80 14.00 0.22 0.00 0.00 0.00 0.00 21.39 42.78 213.90 0.48 0.00 0.00 0.00 3.33 0.00 0.00 0.00 0.53 2.61 5.21 26.05 0.00 5.78 13.22 66.31 0.37 0.04 0.09 0.45 0.01 16.56 33.94 169.74 2.55 34.37 68.74 343.72 5.84 −1.25 −0.87 −4.25 −0.73 0.50 0.50 0.50 0.50 LV bituminous and duckweed biomass

D-5

D-10

−13.18 1.22 0.00 2.40 16.63 2.64 0.00 4.34 0.04 14.09 29.18 −1.01 0.50

−27.23 2.53 0.00 4.80 33.25 5.28 0.00 10.22 0.08 28.93 58.37 −0.51 0.50

D-50

R-1

R-5

R-10

−136.26 −2.35 −14.02 −28.98 12.64 0.22 1.30 2.69 0.00 0.00 0.00 0.00 24.00 2.78 13.89 27.78 166.26 1.01 5.07 10.14 26.39 0.50 2.49 4.97 0.00 0.00 1.58 3.15 51.19 0.39 5.24 12.12 0.42 0.01 0.04 0.09 144.63 2.55 15.58 31.96 291.83 5.84 32.32 64.65 −2.56 −0.74 −1.16 −0.72 0.50 0.50 0.50 0.50 Anhui, China, coal and duckweed biomass

R-50 −145.03 13.45 0.00 138.90 50.69 23.86 15.77 60.74 0.43 159.82 323.24 −3.60 0.50

GHG balances

U-1.6

U-5

U-10

D-1.6

D-5

D-10

U-1.6

U-5

U-10

D-1.6

D-5

D-10

biomass coal butane gasoline diesel kerosene LPG vented CO2 seq CO2 LGHG GHGAF GHGAE GHGI

−1.99 0.51 0.00 6.84 0.00 0.00 0.83 1.05 0.01 7.26 11.00 −0.62 0.70

−6.43 1.64 0.00 21.39 0.00 0.00 2.61 4.11 0.05 23.37 34.37 −0.98 0.70

−12.94 3.30 0.00 42.78 0.00 0.00 5.21 8.46 0.10 46.91 68.74 −1.73 0.70

−1.76 0.45 0.00 0.77 5.32 0.84 0.00 0.38 0.01 6.02 9.34 −0.74 0.70

−5.78 1.48 0.00 2.40 16.63 2.64 0.00 2.14 0.05 19.55 29.18 −1.25 0.70

−11.77 3.00 0.00 4.80 33.25 5.28 0.00 4.92 0.09 39.58 58.37 −1.82 0.70

−2.43 0.62 0.00 6.84 0.00 0.00 0.83 1.48 0.01 7.36 11.00 −0.48 0.70

−7.82 2.00 0.00 21.39 0.00 0.00 2.61 5.32 0.04 23.53 34.37 −0.75 0.70

−15.85 4.05 0.00 42.78 0.00 0.00 5.21 11.20 0.08 47.47 68.74 −0.93 0.70

−2.20 0.56 0.00 0.77 5.32 0.84 0.00 0.92 0.01 6.22 9.34 −0.45 0.70

−7.07 1.81 0.00 2.40 16.63 2.64 0.00 3.62 0.04 20.06 29.18 −0.53 0.70

−14.22 3.63 0.00 4.80 33.25 5.28 0.00 7.19 0.08 40.01 58.37 −1.20 0.70

a

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



Article

GREET model96 and assume transportation distances of 50 miles for feedstocks, 100 miles for products, and 50 miles for CO2. The carbon content of the biomass is used to calculate the amount of CO2 that is removed from the atmosphere in part (e). The GHG emissions avoided from liquid fuels (GHGAF) is calculated by multiplying the total energy of fuels produced by a typical petroleum-based emissions level (i.e., 91.6 kg CO2eq/ GJ LHV), while the GHG emissions avoided from electricity is calculated by multiplying the energy produced by electricity by a typical natural gas-based emissions level (i.e., 101.3 kg CO2eq/ GJ). A negative value of GHGAE represents electricity that is input into the refinery. The GHG emissions index (GHGI) is calculated by dividing the LGHG by the summation of GHGAF and GHGAE. A GHGI value less than 1 represents a process that has superior life-cycle GHG emissions than current processes. As shown in Table 10, the GHGI is equal to 0.50 for all the case studies. This shows that CBTL processes can not only be economically competitive with petroleum-based processes but can achieve lower GHG emissions than current processes. The major contributor to the LGHG is the combustion of liquid fuels. The use of biomass as a feedstock helps mitigate the overall LGHG emissions, thus allowing the refinery to achieve the 50% GHG emissions target. 3.6.2. LV Bituminous Coal and Duckweed Biomass. The amount of carbon from each input and output for each of the six case studies utilizing LV bituminous coal and duckweed biomass is shown in Table 9. For each of these case studies, the amount of carbon originating from coal is about 87% of the total carbon input into the refinery. The total carbon conversion is between 42 and 44% for the unrestricted case studies and 42−45% for the maximum diesel case studies. The GHG balances for these case studies are illustrated in Table 10. The GHGI is equal to 0.70 for all the case studies. By utilizing duckweed biomass together with LV bituminous coal, these CBTL refineries were successful in achieving the 30% reduction in GHG emissions. 3.6.3. Anhui, China, Coal and Duckweed Biomass. Tables 9 and 10 illustrate the carbon and GHG balances, respectively, for the six CBTL refineries that input coal from Anhui, China, and duckweed biomass. For these case studies, the amount of carbon originating from coal is about 84% of the total carbon input into the refinery. The unrestricted case studies achieve a carbon conversion between 45 and 47%, while the case studies that maximized diesel achieved between 45 and 46% carbon conversion. As illustrated in Table 10, the use of duckweed biomass allowed these CBTL refineries to achieve a 30% reduction in GHG emissions compared to petroleum-based processes.

and duckweed biomass across three refinery scales and two different fuel compositions (unrestricted and maximization of diesel). Finally, the last six case studies used coal from the province of Anhui, China, and duckweed biomass across three refinery scales and two different fuel compositions (unrestricted and maximization of diesel). The solution of the MINLP model provided the optimal process topology and was theoretically guaranteed to be between 1 and 10% of the best possible value. For the case studies utilizing LV bituminous coal and hardwood biomass, the overall cost of unrestricted liquid fuels production is $136/bbl for a 1 kBD plant, $84/bbl for a 5 kBD plant, $69/bbl for a 10 kBD plant, and $55/bbl for a 50 kBD plant. These case studies achieved a 50% reduction in GHG emissions compared to current processes. A CBTL refinery inputting LV bituminous coal and duckweed biomass has an overall cost of $122/bbl for a 1.6 kBD plant, $85/bbl for a 5 kBD plant, and $70/bbl for a 10 kBD plant in the unrestricted case. For a CBTL refinery using coal from Anhui, China, and duckweed biomass, the overall costs for a 1.6 kBD plant, a 5 kBD plant, and 10 kBD plant are $117/bbl, $80/bbl, and $64/bbl, respectively. The latter two sets of case studies achieved a 30% reduction in lifecycle GHG emissions of current fossil-fuel processes. As the refinery scale increased past 10 kBD, the effects of economies of scale diminished since several trains were required in the CBTL refinery. A parametric analysis of the coal and biomass price was presented to illustrate the effect of feedstock cost on the break even oil price. The results presented in this paper clearly demonstrate the economic competitiveness of coal and biomass refineries which produce liquid transportation fuels while simultaneously reducing life-cycle greenhouse gas emissions.

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ASSOCIATED CONTENT

S Supporting Information *

The complete mathematical model. This material is available free of charge via the Internet at http://pubs.acs.org.

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AUTHOR INFORMATION

Corresponding Authors

*Tel: (86 10) 62566737. E-mail: [email protected]. *Tel: (609) 258-4595. Fax: (609) 258-0211. E-mail: floudas@ titan.princeton.edu. Notes

The authors declare no competing financial interest.

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

4. CONCLUSIONS A rigorous global optimization framework for the thermochemical production of liquid fuels from coal and biomass has been illustrated in this paper. The framework was able to simultaneously analyze multiple hydrocarbon production and upgrading technologies that produced liquid fuels at prices competitive with petroleum-based processes while achieving a significant reduction in GHG emissions from current processes. Simultaneous heat, power, and water integration was included in the process synthesis framework. The MINLP model was demonstrated on 24 distinct case studies. Twelve case studies utilized a combination of LV bituminous coal and hardwood biomass to produce liquid fuels across four refinery capacities and three different fuel compositions (unrestricted, maximization of diesel, and U.S. ratios). Six case studies input LV bituminous coal

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Coal and Biomass to Liquid Transportation Fuels - American Chemical

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