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Feb 4, 2013 - Biomass and Natural Gas to Liquid Transportation Fuels: Process ... Princeton University, Princeton, New Jersey 08544, United States...
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Biomass and Natural Gas to Liquid Transportation Fuels: Process Synthesis, Global Optimization, and Topology Analysis Richard C. Baliban, Josephine A. Elia, and Christodoulos A. Floudas* Department of Chemical and Biological Engineering, Princeton University, Princeton, New Jersey 08544, United States S Supporting Information *

ABSTRACT: An optimization-based process synthesis framework is proposed for the thermochemical conversion of biomass and natural gas to liquid fuels (BGTL). Hydrocarbons are produced from synthesis gas either directly via Fischer−Tropsch synthesis or indirectly via catalytic conversion of methanol over ZSM-5. Different conversion technologies are investigated to examine their economic effects on BGTL refineries that produce gasoline, diesel, and kerosene. The process synthesis framework includes simultaneous heat, power, and water integration and utilizes a rigorous deterministic global optimization strategy to mathematically guarantee the minimal cost of the BGTL refineries. The refineries have at least 50% less life-cycle CO2 emissions than a standard petroleum-based refinery. Forty-eight case studies are presented to determine the effect of refinery capacity, liquid fuel composition, and biomass feedstock on the overall cost, topological design, material/energy balances, and life-cycle greenhouse gas emissions. Results suggest that these systems can be economically competitive with petroleum-based processes while achieving the 50% emissions reduction.

1. INTRODUCTION

biomass is currently available for biofuel production with an estimated 1.3 billion tons potentially available in the future.6,7 Though biomass feedstocks possess desirable environmental benefits for use as liquid fuel precursors, the overall costs necessary with fuel production can be high due to the cost of the biomass feedstock or the cost needed to convert a cellulosic biomass feedstock. The cost of the biomass feedstock will be highly dependent on the demand for consumption and any opportunity cost associated with land use. Liquid fuels derived from biomass via a thermochemical-based route (i.e., via synthesis gas) require considerable investments of capital to convert the solid raw material to a clean gas product that is devoid of acidic and nitrogenous gas species (e.g., H2S, CO2, NH3). In fact, it has been estimated that the gasification/ cleanup processes needed for biomass plants is approximately 50−65% of the entire plant cost.8,9 The incorporation of additional feedstocks including coal or natural gas has the potential to reduce both the raw material cost and the capital cost for the refinery. Coal/biomass plants have gained interest due to the inexpensive delivered cost of coal ($2.0−$2.5/ million (MM) Btu),3 though the capital cost associated with coal gasification can be higher than biomass gasification due to the high temperatures/pressures needed to operate the gasifier and the extensive cleaning operations necessary to process the raw synthesis gas.10−19 Additionally, the high carbon content of coal may require that a significant portion of the feedstock carbon is converted to CO2, where it is generally vented or sequestered.20−24 If a non-carbon based source of hydrogen was available, the CO2 may be converted to CO via the reverse water gas shift reaction.20,21 However, the current capital and

The transportation sector in the United States currently faces significant challenges over high energy prices, the volatility of the global oil market, and increased pressure to reduce life-cycle greenhouse gas emissions. Geopolitical uncertainty related to the international oil market along with the concern of “peakoil” domestic crude oil production1,2 has motivated national efforts toward greater energy independence. Consequently, the supply demand gap for liquid fuel production between 2010 and 2035 is estimated to be largely satisfied through “nonpetroleum” based feedstocks.3 Though multiple technologies have been proposed to process alternative feedstocks, the U.S. Energy Information Administration projects that a majority of the liquid fuel supply will come from biomass sources.3 Biomass is an attractive feedstock for liquid fuel production and has the potential to play a substantial role in shaping the energy future of the United States. Biomass is a renewable energy source that can absorb atmospheric CO2 during photosynthesis,4−6 so it is capable of addressing both the concerns of enhanced domestic fuel production and life-cycle greenhouse gas emissions reduction. However, the use of land to grow biomass for liquid fuel production must be analyzed in context with other potential land requirements including food, feed, or the preservation of natural habitats. Ultimately, removal of crop or forest residues for biofuel production will have to be investigated using a holistic framework that recognizes the connectivity of multiple processes that occur on both the farm and the ecosystem including soil carbon management, erosion mitigation, nutrient management, water/air quality, and global food/feed/fiber production.6 Nevertheless, it is possible to responsibly develop biomass feedstocks in a sustainable manner to generate a significant national supply without disruption of the human and animal food chain or deforestation.7 In fact, it is estimated that roughly 400−500 million sustainable dry tons of © 2013 American Chemical Society

Received: Revised: Accepted: Published: 3381

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operating costs to produce non-carbon based hydrogen25 generally make this alternative economically unattractive. Recent prospects for shale gas production have helped reduce the delivered cost of natural gas and made this feedstock a more attractive choice for liquid fuel production along with biomass.3,26−29 Natural gas is a very advantageous feedstock for liquid fuel production due to the high hydrogen to carbon ratio within the methane-rich feed. The high ratio can increase the overall yield of carbon in the liquid products, decrease the capital investment required to generate liquid products, and reduce the amount of CO2 that is produced. Hybrid systems cofeeding coal, biomass, and natural gas to produce liquid fuels21−24,30,31 can take advantage of the economic and environmental trade-offs for all three of the major feedstocks. Utilization of the national supplies of coal, biomass, and natural gas in an economic and environmentally suitable fashion is of utmost importance to break the strong dependence that the United States has on petroleum supplies. A recent review has highlighted the key developments made by academic groups worldwide to introduce process design alternatives that can produce gasoline, diesel, and kerosene using any one or a combination of coal, biomass, or natural gas feedstocks.32 The final type of hybrid design involves biomass and natural gas to liquid fuels (BGTL). Previous studies on BGTL processes typically focus on process designs where the topology (i.e., the combination of process units and streams) is fixed. A process simulation is then conducted to determine the heat and mass balances for the process and an economic analysis is performed to determine the viability of the plant.26−29 Though a process design that is developed in this fashion is adequate, this strategy will require a significant amount of computational time and manpower to find the “best possible” design which optimizes some predetermined objective (e.g., maximum profit, minimum cost). Recent developments in process synthesis strategies23,24,30,31,33 have shown that computationally efficient methodologies exist to analyze thousands of process designs simultaneously to extract the optimal process design. A process synthesis strategy has also been developed for switchgrass and natural gas to liquid systems,34 and a proprietary synthesis tool has been developed for gas to liquid (GTL) systems by Shell Global Solutions.35 In this study, an optimization-based process synthesis framework is proposed for directly comparing the technoeconomic and environmental benefits of BGTL processes using a large-scale nonconvex mixed-integer nonlinear optimization (MINLP) model. Several existing or novel process topologies are combined into a superstructure of alternatives for producing the liquid hydrocarbon fuels from biomass and natural gas. Detailed input−output relationships for each process unit in the superstructure are defined for proper unit operation, and the large-scale MINLP model is formulated to encompass the simultaneous consideration of every process unit. A rigorous deterministic global optimization branch-and-bound strategy30 is used to solve the MINLP model to global optimality and theoretically guarantee that the process design selected by the framework will have an objective value that is within a small percentage of the best value possible. Modeling of several components of the process superstructure have been described in detail in previous works,23,24,30,31,33 though all key processes for the BGTL refinery will be discussed in the text below. A simultaneous heat and power integration22,23 using an optimization-based heat-integration approach36 and a series of heat engines is included in the framework to ensure that waste

heat is effectively converted into electricity and process steam utilities.21−23 A comprehensive wastewater treatment network24 that includes a sour stripper, a biological digestor, and a reverse osmosis unit is incorporated using a superstructure approach37−40 to minimize wastewater contaminants and freshwater intake. The process synthesis framework will be utilized to examine (i) biomass gasification with/without recycle synthesis gas, (ii) natural gas conversion via steam reforming or autothermal reforming, (iii) synthesis gas conversion via Fischer−Tropsch (FT) or methanol synthesis, (iv) methanol conversion via methanol to gasoline (MTG) or methanol to olefins (MTO), and (v) hydrocarbon upgrading via ZSM-5 zeolite catalysis, olefin oligomerization, or carbon number fractionation and subsequent treatment. The key products from the BGTL refinery will be gasoline, diesel, and jet fuel (kerosene) with allowable byproducts of liquefied petroleum gas (LPG) and electricity. The case studies in this paper are selected to reflect the quantitative trade-offs associated with different key metrics for the BGTL refinery and are meant to demonstrate the capability of the process synthesis framework. The BGTL process synthesis framework, the large-scale MINLP model, and the global optimization strategies can be readily expanded to incorporate additional refinery technologies or allow for the production of chemical products.

2. BGTL PROCESS SUPERSTRUCTURE: CONCEPTUAL DESIGN AND MATHEMATICAL MODELING This section will outline the design and modeling of the key sections of the BGTL refinery.21,23,24,30,31,33 The complete mathematical model, all relevant nomenclature, and a complete set of process flow diagrams are provided as Supporting Information. 2.1. Biomass Handling and Gasification. The BGTL refinery is assumed to input only one type of biomass feedstock to reduce the complexity of the feedstock processing section and help increase the uniformity of the input to the gasifier. Three distinct categories of feedstocks will be considered (i.e., agricultural residues, perennial grasses, and forest residues), each of which has a representative composition shown in Table 1 that is derived from the ECN Phyllis database.41 The biomass is assumed to be delivered to the BGTL refinery as wood chips (forest residues) or as bales (agricultural residues/perennial feeds) which must be further processed prior to entry to the Table 1. Feedstock Proximate and Ultimate Analysis for Biomass Species41 heating value (kJ/kg)

proximate analysis (db, wt %) feed type

moisture (ar)a

agricultural perennial forest

25 25 45

ash

VMb

FCc

5.1 80.9 14 4.6 79.2 16.2 2.14 N/A N/A ultimate analysis (db, wt

HHVd

LHVe

18 101 18 636 19 130 %)

16 849 17 360 17 842

feed type

C

H

N

Cl

S

O

agricultural perennial forest

46.8 46.9 50.19

5.74 5.85 5.9

0.66 0.58 0.32

0.266 0.501 0

0.11 0.11 0.03

41.4 41.5 41.42

ar, as received. bVM, volatile matter. cFC, fixed carbon. dHHV, higher heating value. eLHV, lower heating value. a

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Figure 1. Biomass gasification flow sheet. Biomass is dried to 20 wt % moisture and then transferred to the gasifier system via a lockhopper. The gasifiers will operate with either a solid fuel (biomass) or a combination of solid and recycle gases as fuel. Residual ash and char that are generated within the gasifiers are separated via the cyclones and recycled to the gasifiers. The raw syngas is then transferred to a tar cracker to remove most of the tar species in the vapor phase.

the heat needed for the reforming reactions and to help crack the tar species in the gasifier. The high temperature of the unit will help to facilitate the water gas shift (WGS) equilibrium of the synthesis gas effluent, though the concentration of the hydrocarbons in the effluent will be far above the equilibrium values. Using the reverse WGS reaction, CO2 may be consumed within the gasifier unit by reaction with H2 that is present within the gasifier. Therefore, any CO2 that is generated by the process can be recycled to the gasifier along with H2 that is produced from pressure-swing adsorption or electrolysis of water. The effluent of the gasifier is passed through a catalytic tar reformer (925 °C), which will reform (i) tar species to CO and H2, (ii) NH3 to N2 and H2, and (iii) C1−C2 hydrocarbons to CO and H2. The current bench-scale performance of a tar reformer from the National Renewable Energy Laboratory has a conversion of 80% of the CH4, 99.6% of tars, 99% of C2H6, 90% of C2H4, and 90% of NH3.42 Equipment is currently being installed for pilot-scale demonstration of the tar reformer performing over a continuous period of time.44 The steam that is present in the gasifier effluent is sufficiently high enough to reform the syngas without the need of additional input steam.44 Heat for the tar reformer is provided by circulating catalyst between the tar reformer and a catalyst regenerator to remove the coke deposits on the catalyst surface. The level of coke deposited on the catalyst is insufficient to provide the heat needed for the endothermic reforming reactions, so additional combustion gases are passed through the regenerator.44 The syngas exiting the tar reformer is cooled to 60 °C and passed to the cleaning section (see Figure 4). 2.2. Natural Gas Conversion. Natural gas is fed to the BGTL refinery at pipeline conditions of 31 bar and 25 °C, and

biomass gasifier. The wood chips are screened to remove particles with size greater than 2 in. and sent to a grinder for further size reduction.42 Any bales of feedstock are sent to the grinder to shred the biomass to a usable size for gasification. Figure 1 details the process flow diagram for the generation of synthesis gas from biomass. The moisture content of the biomass must be reduced to 20 wt % through a preliminary drying step43 before the feedstock can be injected to the gasifier. Any heat necessary for the dryer is provided by flue gas generated through combustion within the refinery. The flue gas leaves the dryer at 110 °C and 1.05 bar, and is passed through an air cyclone and a baghouse filter to remove any particulates that are present.42 The heated biomass leaves the dryer at 105 °C and 1.05 bar, and is transferred to a high-pressure gasifier (30 bar) using compressed CO2 (10 wt %) and a lockhopper. Readers are referred to the studies by Baliban et al.21,23 for the detailed breakdown of the biomass species in the gasifier. The effluent of the biomass gasifier is a mixture of synthesis gas, C1−C2 hydrocarbons, acid gases (e.g., NH3, H2S), tar, char, and ash.21,23 The solid ash and char are separated from the vapor phase using cyclones, and are recycled back to the gasifier. It is assumed that the recycle of char will effectively provide a 100% conversion of the carbon in the biomass to vapor species while the ash is removed from the gasifier as slag. The composition of the gasifier effluent is a function of the biomass composition, gasifier temperature, and oxidizer flow rate, and a detailed mathematical model has been formulated to determine the composition21,23 (see Supporting Information). The gasifier is assumed to operate between 800 and 1000 °C and will utilize steam for gasification of biomass/char, reforming of C1−C2 hydrocarbons, and reforming of tar species. High-purity oxygen is input to the gasifier to provide 3383

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Figure 2. Natural gas conversion flow sheet. Natural gas is combined with recycle methane and may be converted to synthesis gas (CO, CO2, H2, H2O) via steam reforming or autothermal reforming.

Figure 3. Natural gas utilities flow sheet. Natural gas and recycle fuel gas may be utilized to produce electricity through a gas turbine or additional process heat via a fuel combustor. The effluents from both of these processes are cooled and then are either vented or passed over a CO2 recovery unit to capture and process the produced CO2.

the NETL Quality Guidelines for Energy Systems Studies report and is based on the mean of over 6800 samples of pipeline

is utilized in one of six major processes (see Figures 2 and 3). The input natural gas composition (see Table 2) is taken from 3384

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rate, so the species molar flow rates (NS) are sufficient to accurately define the equilibrium constraint. The SMR equilibrium constraint utilizes molar species concentrations (xS) to account for the change in total molar flow rate. The equilibrium constant in eq 2 was adjusted from the value extracted from Aspen Plus v7.3 (KMR,o SMR ) for the higher pressure of the reforming unit using eq 3 (PSMR = 30 bar).

Table 2. Molar Compositions (x) of All Species in the Input Natural Gas species

x

species

x

species

x

CH4 CO2

0.931 0.010

C2H6 C3H8

0.032 0.007

N2 n-C4H10

0.016 0.004

quality natural gas.45,46 Natural gas can be desulfurized using a zinc oxide polishing bed (sulfur guard) to clean any mercaptanbased odorizers from the gas to prevent catalyst contamination with the reformers.46 Natural gas and other methane-rich recycle gases may also be sent to a gas turbine to produce electricity or to a fuel combustor to provide process heat (see Figure 3).23 CO2 produced from these units may be captured using flue-gas capture technology and mixed with additional process CO2 for appropriate handling (see Figure 4). 2.2.1. Steam Reforming. Steam reforming (or autothermal reforming) involves indirect conversion of the natural gas to synthesis gas (syngas; CO, CO2, H2, H2O; Figure 2). Steam reforming of the natural gas uses a nickel-based catalyst contained inside high alloy steel tubes. Heat is provided for the endothermic reforming of methane via combustion of recycle fuel gas and additional input natural gas over the outside of the tubes. The reformer operates at a pressure of 30 bar with typical reaction temperatures of 800−900 °C.46 The effluent reformed gas will be constrained by both WGS equilibrium (eq 1) and steam methane reforming (SMR) equilibrium (eq 2). The effluent concentrations of C2 and higher hydrocarbons are assumed to be negligible with respect to the concentration of methane. The WGS equilibrium conserves the total molar flow

S S WGS,o S S NSMR, u ,H 2ONSMR, u ,CO = K SMR NSMR, u ,H 2NSMR, u ,CO2

(1)

3

S S MR S S xSMR, u ,CH4xSMR, u ,H 2O = K SMR xSMR, u ,H 2 xSMR, u ,CO MR KSMR

=



(2)

⎞2

MR,o 1.01325 KSMR ⎜ ⎟ ⎝ PSMR ⎠

(3)

The equilibrium constants provided by Aspen for a pressure of 1 atm (KWGS,o and KMR,o) are shown in Table 3 for 800−1000 Table 3. Chemical Equilibrium Constants (KMR,o and KWGS,o) Extracted from Aspen Plus v7.3 for Steam Methane Reforming (SMR) and Water Gas Shift (WGS)a

a

temp (°C)

SMR

WGS

800 850 900 950 1000

74.165 228.599 640.874 1652.908 3960.340

1.08152 0.91372 0.78487 0.68400 0.60368

The values shown are for a pressure of 1 atm.

Figure 4. Synthesis gas (syngas) handling flow sheet. Syngas may be passed over a forward/reverse water gas shift reactor to alter the H2 to CO/CO2 ratio prior to Fischer−Tropsch or methanol synthesis. The syngas is then cooled, flashed to remove water, and may be directed to an acid-gas removal system for CO2 and H2 removal. The captured CO2 may be vented, sequestered, or recycled back to process units. 3385

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°C. Note that the equilibrium constant does not need to be adjusted for the WGS equilibrium due to conservation of moles. The steam reformer operating temperature will be selected from a discrete set of three options, namely 800, 850, or 900 °C. By studying alternative modeling methods for defining the constraints in the steam reformer (e.g., molar species concentrations in WGS equilibrium), the current mathematical formulation is selected for the steam reformer since it provided the best computational performance for this study. All nonhydrocarbon and nonsyngas species (e.g., N2, Ar) are assumed to be inert. The effluent reformed gas is directed to syngas cleaning (see Figure 4). Ambient air (13 °C, 1.01 bar) is compressed to 1.1 bar to provide a 20 mol % stoichiometric excess of oxygen needed for combustion of the fuel gas within the reformer. The combusted fuel gas exits the reformer at 640 °C, is cooled to 120 °C to recover waste heat, and is then directed to either the stack or a CO2 recovery unit. 2.2.2. Autothermal Reforming. Autothermal reforming of the natural gas will input a combination of steam for endothermic reforming and high-purity oxygen for partial combustion within the same reactor. The autothermal reformer (ATR) will operate at a pressure of 30 bar with a temperature between 900 and 1000 °C. Oxygen is provided through cryogenic air separation (99.5 wt %) or electrolysis of water (100 wt %), and is preheated to 300 °C prior to entering the reformer. Steam will also be preheated to 550 °C and the natural gas will be preheated to 300 °C to reduce the oxygen requirement within the reformer. The molar ratio of steam to total carbon entering the reformer will vary between 0.5 and 1.5, and the effluent will be governed by the WGS equilibrium (eq 4) and SMR equilibrium (eq 5). The choice of mathematical formulation of the autothermal effluent is similar to that of the steam reformer and is based on computational performance. The equilibrium constants for the autothermal reformer are shown in Table 3 for the three possible operating temperatures of the reactor (i.e., 900, 950, or 1000 °C). The constant for steam methane reforming is adjusted for a 30 bar (PATR = 30 bar) reactor using eq 6. S S WGS,o S S N ATR, u ,H 2ON ATR, u ,CO = KATR N ATR, u ,H 2N ATR, u ,CO2

(u ′ ,POM,CH4) ∈ S

∀s∈

S NOCO, u , s = fcOCO, s

(6)

∑ UF

(u ′ ,POM,CH4) ∈ S

NuS ,POM,CH4 ′



NuS ,OCO, s ′ UF

HC ∀ s ∈ SOCO

(9) S NOCO, u , s AR s ,C = cd OCO, s

∑ HC s ∈ SOCO

Ef SOCO

∑ (u ′ ,OCO, s) ∈ S

UF

NuS ,OCO, s ′ (10)

2.3. Synthesis Gas Cleaning. The process flow diagram for processing the raw syngas is shown in Figure 4. The syngas effluent from the reformers may be directed to a separate water gas shift reactor that operates without the presense of sulfur species. The unit will operate at a pressure of 28 bar and a temperature between 400 and 600 °C. Alternatively, the reformer effluent may bypass the water gas shift reactor or combine with the biomass syngas entering the sour water gas shift unit. This latter option may be beneficial to reduce the capital costs associated with the gas shift reactor, though it may increase the costs of the CO2/H2S acid gas separation system. Sulfur-free syngas will be cooled to 35 °C to remove wastewater via vapor−liquid equilibrium, and may be directed to an acid gas recovery unit for CO2 removal. Alternatively, the sulfur-free syngas may be directly sent to the hydrocarbon production section and consume CO2 via the reverse water gas shift reaction. The syngas from the sour water gas shift unit is combined with the syngas exiting the tar cracker, cooled to 185 °C, and then sent to a scrubbing system (SCRUB) to remove residual tars, particulates, and NH3 as part of a wastewater stream. The

The effluent from the autothermal reformer is directed to the synthesis gas cleaning section (see Figure 4). 2.2.3. Direct Conversion to Methanol via Partial Oxidation. Natural gas may be directly converted to methanol via gas phase partial oxidation operated by a free radical mechanism.33,47−49 The natural gas is compressed to 52 bar and then passed into a quartz-lined tubular reactor (POM) operating at 450 °C and 50 bar.47 The per-pass conversion of methane (fc) is 13%49 (eq 7) with a carbon distribution (cd) of 63% to CH3OH, 30% to CO, 6% to CO2, and 1% to C2H647 (eq 8), where SEf POM represents the set of species that are formed from conversion of the methane. S NPOM, u ,CH4 = fcPOM,CH4

(8)

(u ′ ,OCO, s) ∈ S

∀s∈

⎛ ⎞2 MR MR,o 1.01325 = KATR KATR ⎜ ⎟ ⎝ PATR ⎠

NuS ,POM,CH4 ′

Under the reaction conditions assumed in this study, all formaldehyde is assumed to decompose quickly to H2 and CO.47 The effluent from the reactor is combined with the effluent from the methanol generated from synthesis gas, cooled to 35 °C, and flashed to separate the methanol/water mixture. The recycle gases are either (i) recompressed and recycled to the POM reactor, (ii) heated to 500 °C and expanded to 30 bar for use in a gas turbine, or (iii) heated to 500 °C and expanded to 1.3 bar for use as fuel gas. The crude methanol/water mixture is combined with additional methanol from the plant prior to degassing and subsequent processing. 2.2.4. Direct Conversion to Olefins via Oxidative Coupling. Natural gas can be contacted with a reducible metal oxide catalyst to promote oxidative dehydrogenation via free radical formation.33,50−58 The reactor (OCO) is assumed to operate at 800 °C and 3.8 bar55 with suitable expansion of the natural gas from the pipeline pressure to recover electricity from a turbine. This study will assume that the per-pass conversion of CH4 is 25% (eq 9) with the product distributed between C2−C7 hydrocarbons, CO, CO 2 , and coke.33 The catalyst is regenerated (OCO-CAT) by passing air (10% stoichiometric excess of O2) over the catalyst surface for reoxidation and removal of the coke to CO2. The flue gas is cooled to 120 °C to recover waste heat and is either vented or sent to a CO2 recovery unit. The effluent of the reactor is cooled to 35 °C for water knockout (OCO-F), compressed to 50 bar, and then sent to a CO2 removal unit (OCO-CO2). The effluent from the CO2 removal unit is then directed to the MOGD reactor to generate gasoline and distillate.

3

(5)

UF

Ef SPOM

(4)

S S MR S S xATR, u ,CH4xATR, u ,H 2O = KATR xATR, u ,H 2 xATR, u ,CO



S NPOM, u , s AR s ,C = cdPOM,CH4

(7) 3386

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Figure 5. Acid gas removal system flow sheet. Acid gases may be directed to one of three absorption systems to remove the CO2 and H2S and produce a clean gas. The H2S gases are then directed to either a Claus plant or a LO-CAT system for recovery of the sulfur. For clarity, the sulfurrich acid gases are shown as dotted lines, the tail gas is shown as a dashed line, and the clean gases are shown as orange lines.

wastewater is sent to a biological digestor (Supporting Information, Figure S12) to treat the organic contaminants while the scrubbed syngas is passed over an acid gas removal (AGR) system to co-remove the H2S and CO2 from the syngas. The AGR system is shown in Figure 5 and consists of a singlecapture amine separation system,46 a single-capture methanol absorption system,43 or a dual-capture methanol absorption system.8 The single-capture systems will cocapture H2S and CO2 in one stream, which must then be directed to a sulfur recovery unit (SRU). The dual-capture system will capture both H2S and CO2 but in separate streams to allow for separate downstream process handling. Note that the acid gas unit must be utilized in the process to remove the sulfur species and prevent poisoning of the downstream hydrocarbon production catalysts. All of the capture units will remove 100% of the H2S and 90% of the CO2 from the input gases. Approximately 3 mol of CO2/mol of H2S is entrained with the H2S in the dualcapture system.8 The heavy steam or electricity requirements necessary to perform the absorption/deabsorption process are considered in the process synthesis model. The clean gases are assumed to leave the capture units at 25 bar and 35 °C, while the acid gases will leave at 1 bar and 35 °C. The CO2 from the dual-capture acid gas recovery unit may be (a) compressed to 31 bar for recycle to the reformers or the water gas shift units or (b) compressed to 150 bar for sequestration. Note that both compression options will utilize multiple compression stages with intercooling to control the temperature rise. The CO2 may alternatively be vented to the atmosphere. The SRU in the BGTL refinery will either be a Claus recovery system19 or a LO-CAT iron chelate based process.42

The Claus process will contain an oxygen-blown furnace that can convert up to 95% of the H2S to solid sulfur. The sulfur in the tail gas will be present in various forms (e.g., H2S, SO2, COS, S) which can be hydrogenated to H2S. To recover the sulfur in the tail gas, the tail gas may be recycled back to one of the raw syngas acid gas recovery units. Alternatively, the tail gas may be sent to a dedicated amine recovery unit, as is practiced by Shell in the Shell Claus Off-gas Treating (SCOT) process. If the LO-CAT process is utilized, essentially a 100% recovery of the H2S to sulfur can be realized with a single pass through the system.42 2.4. Hydrocarbon Production/Upgrading. 2.4.1. Fischer−Tropsch Hydrocarbon Production. The hydrocarbon production section (Supporting Information, Figures S6 and S9) will convert the syngas using either FT synthesis or methanol synthesis. The FT units will operate at 20 bar and will utilize either a cobalt-based or iron-based catalyst.23,24,31 Cobalt-based catalysts will not facilitate the water gas shift reaction, so the FT units will be ideal for achieving high perpass conversion of the CO to FT liquids. However, the extent of catalyst oxidative degration due to high water partial pressure is a contentious topic, and current data imply that the effect of catalyst stability based on the presense of water is not clear.9 Though cobalt oxidation is reversible, increased time on stream will result in higher levels of methane formation and lower levels of C5+ liquids from the FT unit.9 For this study, the CO per-pass conversion was set to 60% to avoid catalyst oxidation, though conversion levels may approach 80% if catalyst stability can be achieved.9 Note that both high and low temperature cobalt-based FT systems are considered in this analysis. To 3387

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three-phase separator to remove the aqueous phase from the residual vapor and any hydrocarbon liquid. Any oxygenates that are present in the vapor phase may be removed using an additional separation unit. The water lean FT hydrocarbons are then sent to a hydrocarbon recovery column for fractionation and further processing (Supporting Information, Figure S8). The oxygenates and water removed from the stream are mixed and sent to the biological digestor for wastewater treatment. The FT hydrocarbons may also be passed over a ZSM-5 catalytic reactor operating at 408 °C and 16 bar59 to be converted into mostly gasoline range hydrocarbons and some distillate.59,60 The ZSM-5 unit will be able to convert the oxygenates to additional hydrocarbons, so no separate processing of the oxygenates will be required for the aqueous effluent. The raw product from FT-ZSM5 is fractionated to separate the water and distillate from the gasoline product. The water is mixed with other wastewater knockout, and the distillate is hydrotreated to form a diesel product. The raw ZSM-5 HC product is sent to the LPG−gasoline separation section for further processing (Supporting Information, Figure S10). The water lean FT hydrocarbons are sent to a hydrocarbon recovery column, as shown in Figure S8 in the Supporting Information. The hydrocarbons are split into C3−C5 gases, naphtha, kerosene, distillate, wax, offgas, and wastewater.21,62 The upgrading of each stream will follow a detailed Bechtel design62,63 which includes 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. 2.4.3. Methanol Synthesis. The methanol synthesis reactor (Supporting Information, Figure S9) will operate at 300 °C and 50 bar and will input a sulfur-free synthesis gas that contains a Ribblett ratio equal to 1. The syngas leaving the cleaning section must be compressed to 51 bar prior to entering the methanol synthesis reactor. The methanol synthesis reactor will assume that equilibrium is achieved for the methanol synthesis reaction (eq 11) and the water gas shift reaction (eq 12).

date, only the low-temperature FT unit has been commercially available, but this study investigates the possibility of a hightemperature system using an α value that is consistent with high-temperature FT operation (e.g., α = 0.72). The iron-based catalysts will operate using either low or high temperature and will facilitate equilibrium of the water gas shift reaction within the FT units. Therefore, these reactors could consume CO2 within the unit using H2 to produce the CO necessary for the FT reaction.23,24,31 However, consumption of CO2 in these units will not be possible unless the CO2/(CO + CO2) ratio is above some critical threshold. Due to the high equilibrium constant for the water gas shift reaction at FT temperatures, it is strongly preferable for CO to be converted to CO2 unless the initial charge of CO2 is high enough to force the reverse reaction. In addition to the CO2/(CO + CO2) ratio, the amount of inlet hydrogen with respect to both CO and CO2 will play an important role in FT operation. Low-temperature iron-based units have been operated successfully with inlet H2/CO ratios between 0.5 and 19,59,60 due to in situ water gas shift activity that effectively produces an appropriate outlet H2/CO ratio near 1.7−2.0. Though these FT processes require substantially less hydrogen than FT processes with a 2/1 inlet ratio, approximately 50% of the CO is converted to CO2. To prevent such a large increase in the outlet CO2 concentration, the inlet ratio of H2/(CO + CO2) must be set to ensure that CO2 can be used as a carbon source due to the reverse water gas shift reaction. The Ribblett ratio9,61 is defined such that H2/(2CO + 3CO2) is approximately equal to 1, and is highly useful because the effluent composition of unreacted syngas from an FT unit will maintain roughly the same value as the inlet. Therefore, the internal or external gas loop designs for FT synthesis can be theoretically designed such that very high conversion rates of CO and CO2 are achieved in the BGTL refinery. To examine the effects of the H2/(CO + CO2) ratio, the synthesis gas entering the iron-based FT units will handled in one of two ways. One low-temperature unit (240 °C) and one high-temperature unit (320 °C) will require an inlet Ribblett ratio that is equal to 1, and should help to facilitate the reverse water gas shift (rWGS) reaction as the CO2 inlet concentration increases. The other two units will operate at 267 °C and have an effluent composition that is based on two previous DOE reports.31,59,60 These two units will have an H2/CO inlet ratio between 0.5 and 0.7 and will ensure that the H2/CO ratio in the effluent is equal to 1.7 from forward water gas shift (fWGS) conversion. Hydrogen may be recycled to any of the FT units to shift either the H2/CO ratio or the H2/CO2 ratio to the appropriate level. Steam may alternatively be used as a feed for the two iron-based fWGS FT units to shift the H2/CO ratio in situ. The two streams exiting the cobalt or iron FT units will be a waxy liquid phase and a vapor phase containing a range of hydrocarbons. The wax will be directed to a hydrocracker (WHC), while the vapor phase is split (SPFTH) for further processing. 2.4.2. Fischer−Tropsch Hydrocarbon Upgrading. The vapor phase effluent from FT synthesis will contain a mixture of C1−C30+ hydrocarbons, water, and some oxygenated species. Figure S7 in the Supporting Information details the process flow sheet used to process this effluent stream. The stream will be split and can pass through a series of treatment units designed to cool the stream and knock out the water and oxygenates for treatment. Initially, the water-soluble oxygenates are stripped from the stream. The stream is then passed to a

CO + 2H 2 ↔ CH3OH

(11)

CO2 + H 2 ↔ CO + H 2O

(12)

The concentration of CO2 input to the methanol synthesis reactor will have significant ramifications on the downstream processing in the BGTL refinery. If the inlet CO2 composition is low due to the use of precombustion capture, then the extent of the water gas shift reaction will be negligible and the concentration of both CO2 and H2O in the effluent stream will be minimal. This ultimately allows for a less energy intensive methanol purification if it is required for downstream conversion. Additionally, the conversion of CO to methanol can approach 45−50% depending on the concentration of inert species that are present in the synthesis reactor. As the inlet concentration of CO2 increases, the reverse water gas shift reaction will begin to occur with an increased production of H2O and a decreased per-pass conversion of CO + CO2. The presence of H2O along with methanol will increase the cost of methanol purification, but if the methanol convesion units are tolerant to high water concentrations, then the purification step will not be needed. In fact, operation of the methanol synthesis reactor with a Ribblett ratio of 1 allows for a very high overall conversion of CO + CO2 using subsequent recycle of the unreacted syngas. Unlike FT synthesis, no light hydrocarbon 3388

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CH4, 6.5 wt % C2−C4 paraffins, 56.4 wt % C2−C4 olefins, and 35.7 wt % C5−C11 gasoline.64 The MTO unit is modeled mathematically using an atom balance and a typical composition seen in the literature.64 The MTO product is fractionated (MTO-F) to separate the light gases, olefins, and gasoline fractions. The MTO-F unit is assumed to operate as a separator unit where 100% of the C1−C3 paraffins are recycled back to the refinery, 100% of the C4 paraffins and 100% of the olefins are directed to the MOGD unit, 100% of the gasoline is combined with the remainder of the gasoline generated in the process, and 100% of the water generated in the MTO unit is sent for wastewater treatment. The separated olefins are sent to the MOGD unit, where a fixed bed reactor is used to convert the olefins to gasoline and distillate over a ZSM-5 catalyst. The gasoline/distillate product ratios can range from 0.12 to >100, and the ratio chosen in this study was 0.12 to maximize the production of diesel. The MOGD unit operates at 400 °C and 1 bar and will utilize steam generation to remove the exothermic heat of reaction within the unit. The MOGD unit is modeled with an atom balance and will produce 82% distillate, 15% gasoline, and 3% light gases.64 The product will be fractionated (MTODF) to remove diesel and kerosene cuts from the gasoline and light gases. The MTODF unit will be modeled as a separator unit where 100% of the C11−C13 species are directed to the kerosene cut and 100% of the C14+ species are directed to the diesel cut. 2.4.5. LPG−Gasoline Separation. The LPG and gasoline generated from ZSM-5 conversion of the FT hydrocarbons or the methanol must be passed through a series of separation units to extract the LPG from the gasoline and alkylate any isobutane to a blending stock for the final gasoline pool (Supporting Information, Figure S10). Light gases are initially removed via one of two knockout units, and the crude hydrocarbons are passed through a de-ethanizer column, a stabilizer column, an absorber column, a splitter column, and an LPG alkylate splitter to separate the LPG from the gasoline fractions. Each of these units is modeled mathematically as a splitter unit where the split fraction of each species to an output stream is given by the information in the Process Flow Diagrams P850-A1501 and P850-A1502 from the NREL study.42 All low pressure steam and cooling water needed for each of the units are derived for each of the units in the NREL study. The total amount of process utility that is needed per unit flow rate from the top or bottom of the column is calculated, and this ratio is used as a parameter in the process synthesis model to determine the actual amount of each utility needed based on the unit flow rate. In addition to the distillation columns within this section, there is also an alkylation unit that is used to convert isobutane and butene to an alkylate blending stock for the gasoline pool. The alkylate was modeled as isobutane,42 and the alkylation unit was modeled using a species balance where the key species, butene, was completely converted to isobutane. Butene is used as the limiting species in this reaction because it is generally present in a far smaller concentration than isobutane. 2.5. Hydrogen/Oxygen Production. Hydrogen is produced via pressure-swing adsorption or an electrolyzer unit, while oxygen can be provided by the electrolyzer or a separate air separation unit (Supporting Information, Figure S11). 2.6. Wastewater Treatment. A complete wastewater treatment network (Supporting Information, Figures S12 and S13) is incorporated that will treat and recycle wastewater from various process units, blowdown from the cooling tower,

gases will be formed during methanol synthesis, so an internal recycle can be employed after appropriate recompression to the feed inlet pressure. For this study, it is assumed that the methanol conversion units (ZSM-5 based reactors) can tolerate crude methanol containing up to 50 wt % water. Therefore, no additional methanol purification step is required between methanol synthesis and methanol conversion. Note that high levels of water in the crude methanol stream are not anticipated to be a concern because the downstream methanol processing units will yield 50 wt % water from the hydrocarbon synthesis.64 The effluent from the methanol synthesis reactor is cooled to 35 °C, and a crude methanol stream is separated using vapor− liquid equilibrium at 48 bar. The amount of methanol that is entrained in the vapor phase is dependent on the input concentration of syngas to the flash unit, but a majority (over 95%) of the methanol can be recovered by enforcing a stoichiometric amount of H2 in the input to the synthesis reactor (i.e., Ribblett ratio = 1). The vapor stream from the flash unit is split so that 5% may be purged to remove inert species and the remaining 95% is compressed to 51 bar and then recycled to the methanol synthesis reactor. The purge stream is recycled back to the process and used as fuel gas. The crude methanol product from the flash unit is heated to 200 °C, expanded to 5 bar to recover electricity, and then cooled to 60 °C prior to entering a degasser distillation column. The degasser will remove all of the entrained gases from the liquid methanol/water while recovering 99.9% of the methanol. The entrained gases are recycled back to the process for use as fuel gas. The bottoms from the degasser will contain methanol and water, with a methanol composition dependent on the level of CO2 input to the synthesis unit. 2.4.4. Methanol Conversion. The crude methanol is split to either the methanol-to-gasoline (MTG) process or to the methanol-to-olefins (MTO) and Mobil olefins-to-gasoline/ distillate (MOGD) processes. The MTG process will catalytically convert the methanol to gasoline range hydrocarbons using a ZSM-5 zeolite and a fluidized bed reactor. The MTG effluent is outlined in Table 3.4.2 of the Mobil study65 and in Process Flow Diagram P850-A1402 of the NREL study.42 Due to the high level of component detail provided by NREL for both the MTG unit and the subsequent gasoline product separation units, the composition of the MTG reactor used in this study is based on the NREL report. The MTG unit will operate adiabatically at a temperature of 400 °C and 12.8 bar. The methanol feed will be pumped to 14.5 bar and heated to 330 °C for input to the reactor. The methanol will be converted to 44 wt % water and 56 wt % crude hydrocarbons, of which 2 wt % will be light gas, 19 wt % will be C3−C4 gases, and 19 wt % will be C5+ gasoline.42 The crude hydrocarbons, upon passing through the upgrading steps detailed in the subsequent sections, will ultimately be separated into finished fuel products, of which 82 wt % will be gasoline, 10 wt % will be LPG, and the balance will be recycle gases. This is modeled mathematically in the process synthesis model by using an atom balance around the MTG unit and assuming a 100% conversion of the methanol entering the MTG reactor.42,65 Any methanol entering the MTO process unit is heated to 400 °C at 1.2 bar. The MTO fluidized bed reactor operates at a temperature of 482 °C and a pressure of 1 bar. The exothermic heat of reaction within the MTO unit is controlled through the generation of low-pressure steam. One hundred percent of the input methanol is converted into olefins containing 1.4 wt % 3389

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Table 4. BGTL Refinery Wastewater Treatment Reference Capacities, Costs (2011 Dollars), and Scaling Factors

a

description

Co (MM $)

So

SMax

unitsa

scale basis

sf

ref

biomass handling (forest) biomass handling (nonforest) biomass gasification, tar cracking, and gas cooling autothermal reformer steam methane reformer water gas shift single-capture amine absorption single-capture methanol absorption dual-capture methanol absorption Claus plant LO-CAT system

4.65 14.42 55.22 10.26 63.74 3.75 68.7 32.1 58.3 27.6 4.2

17.9 17.9 17.9 12.2 26.1 107.9 62.37 54.9 54.9 1.59 52.2

30.6 30.6 33.3 35.0 35.0 150.0b 100.0b 192.0 192.0 10.0 68.0

kg/s kg/s kg/s kg/s kg/s kg/s kg/s kg/s kg/s kg/s kg/h

as-received biomass as-received biomass dry biomass natural gas feed natural gas feed feed gas feed gas feed gas feed gas recovered sulfur recovered sulfur

0.77 0.77 0.67b 0.67b 0.67b 0.67b 0.63 0.63 0.63 0.67b 0.65

8 8 43 62 46 46 46 8 8 46 42

For So and SMax. bMaximum capacity (SMax) and scaling factor (sf) are estimated.

blowdown from the boilers, and input freshwater.24 Process wastewater is treated using a sour stripper or a biological digestor to remove the sulfur species (e.g., H2S), nitrogen species (e.g., NH3), or hydrocarbon species (e.g., oxygenates) that are expected to be in the wastewater streams. Clean output of the network includes (i) process water to the electrolyzers, (ii) steam to the gasifier, autothermal reformer, steam reformer, and water gas shift reactor, and (iii) discharged wastewater to the environment. 2.7. Unit Costs. The total direct costs, TDC, for the BGTL refinery hydrocarbon production and upgrading units are calculated using estimates from several literature sources19,42,59,60,65 using the cost parameters in Table 4 and eq 13 TDP = (1 + BOP)Co

S sf So

Cost uU =

(15)

The levelized costs for the units described for natural gas conversion are added to the complete list of BGTL process units in previous studies.23,24,30,31,33 2.8. Objective Function. The objective function for the model is given by eq 16. The summation represents the total cost of liquid fuel production and includes contributions from the feedstock cost (CostF), the electricity cost (CostEl), the CO2 sequestration cost (CostSeq), and the levelized unit investment cost (CostU). Each of the terms in eq 16 is normalized to the total volume of products produced (Prod). Note that other normalization factors (e.g., total volume of gasoline equivalent, total energy of products) and other objective functions (e.g., maximizing the net present value) can be easily incorporated into the model framework.

(13)

where Co is the installed unit cost, So is the base capacity, S is the actual capacity, sf is the cost scaling factor, and BOP is the balance of plant (BOP) percentage (site preparation, utility plants, etc.). The BOP is estimated to be 20% of the total installed unit cost. All capital cost numbers are converted to 2011 dollars using the Chemical Engineering Plant Cost Index.66 The cost estimates for the biomass handling and gasification units, the two natural gas conversion technologies, and the acid gas removal technologies are included in Table 4. Cost estimates for all other process units in the BGTL refinery are taken from various literature sources8,19,42,43,46,62 and previous works.23,24,31,33 The total plant cost, TPC, for each unit is calculated as the sum of the total direct capital, TDC, plus the indirect costs, IC. The IC include engineering, startup, spares, royalties, and contingencies and is estimated to be 32% of the TDC. The TPC for each unit must be converted to a levelized cost to compare with the variable feedstock and operational costs for the process. Using the methodology of Kreutz et al.,8 the capital charges (CC) for the refinery are calculated by multiplying the levelized capital charge rate (LCCR) and the interest during construction factor (IDCF) by the total overnight capital (eq 14). CC = LCCR· IDCF· TPC

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

min CostFBio + CostFNG + Cost FH2O + CostFBut + CostEl + CostSeq +

∑ u ∈ UInv

Cost uU (16)

The process synthesis model with simultaneous heat, power, and water integration represents a large-scale nonconvex mixedinteger nonlinear optimization (MINLP) model that was solved to global optimality using a branch-and-bound global optimization framework.30 At each node in the branch-andbound tree, a mixed-integer linear relaxation of the mathematical model is solved using CPLEX67 and then the node is branched to create two children nodes. The linear relaxation was constructed using a piecewise-linear partitioning scheme that depends logarithmically on the number of binary variables for all nonlinear terms. The solution pool feature of CPLEX is utilized during the solution of the relaxed model to generate a set of distinct points (150 for the root node and 10 for all other nodes), each of which is used as a candidate starting point to solve the original model. For each starting point, the current binary variable values are fixed and the resulting NLP is minimized using CONOPT.68 If the solution to the NLP is less than the current upper bound, then the upper bound is replaced with the NLP solution value. At each step, all nodes that have a lower bound that is within an ε tolerance of the current upper bound (LBnode/UB ≥ 1 − ε) are eliminated from the tree. For a more complete coverage of branch-and-bound algorithms, the reader is directed to the textbooks of Floudas69,70 and reviews of global optimization methods.71−73

(14)

8

Kreutz et al. calculates an LCCR value of 14.38%/year and IDCF of 7.6%. Thus, a multiplier of 15.41%/year is used to convert the TPC into a capital charge rate. Assuming an operating capacity (CAP) of 330 days/year and operation/ maintenance (OM) costs equal to 5% of the TPC, the total levelized cost (CostU) associated with a unit is given by eq 15. 3390

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3. COMPUTATIONAL STUDIES Forty-eight distinct case studies were used to examine the capability of the process synthesis model. The natural gas feedstock composition was based on an average representation across multiple well heads (Table 2), while the biomass was selected from an average representation of either an agricultural residue, a forest residue, or a perennial crop (Table 1). The global optimization framework was terminated if all nodes in the branch-and-bound tree were processed or if 100 CPU hours had passed.30 Four representative capcities of 1, 5, 10, and 50 thousand barrels per day (kBD) were chosen to examine the effects of economy of scale on the BGTL refinery. The liquid fuel production of gasoline, diesel, and kerosene was selected to either (a) represent the 2010 demand in the United States (i.e., 67 vol % gasoline, 22 vol % diesel, 11 vol % kerosene),74 (b) maximize the diesel production (i.e., ≥75 vol %), (c) maximize the kerosene production (i.e., ≥70 vol %), or (d) freely output any unrestricted composition of the products. These case studies will be labeled as N-C, where N represents the type of product composition (i.e., R, 2010 U.S. ratios, D, maximum diesel, K, maximum kerosene, U, unrestricted composition) and C represents the capacity in kBD. For example, the U-1 label represents the 1 kBD capacity refinery with an unrestricted product composition. All of the case studies will ensure that the life-cycle GHG emissions from the refinery are at least 50% less than the emissions of current fossil fuel based processes. Thus, the lifecycle GHG emissions must be at most 50% of a petroleumbased refinery (91.6 kg of CO2eq/GJLHV)75 for the liquid fuels or that of a natural gas combined cycle (NGCC) plant (101.3 kg of CO2eq/GJ) for electricity.19 Note that the emissions for the natural gas based electricity production process is calculated from the NGCC case without carbon capture and sequestration in the NETL report.19 If electricity is input to the BGTL refinery, then the associated GHG emissions with electricity production are added to the life-cycle GHG emissions for the refinery. If electricity is output from the BGTL refinery, the avoided GHG emissions are subtracted from the total life-cycle GHG emissions. The cost parameters used for the BGTL refinery are listed in Table 5. The costs for feedstocks (i.e., biomass, natural gas,

transportation, storage, and monitoring of the CO2 is shown in Table 5. Once the global optimization algorithm has completed, the resulting process topology provides (i) the operating conditions and working fluid flow rates of the heat engines, (ii) the amount of electricity produced by the engines, (iii) the amount of cooling water needed for the engines, and (iv) the location of the pinch points denoting the distinct subnetworks. Given this information, the minimum number of heat exchanger matches necessary to meet specifications (i)−(iv) are calculated as previously described.23,24,69,76 Upon solution of the minimum matches model, the heat exchanger topology with the minimum annualized cost can be found using the superstructure methodology.22,69,76 The investment cost of the heat exchangers is added to the investment cost calculated within the process synthesis model to obtain the final investment cost for the superstructure. 3.1. Optimal Process Topologies. The major topological selections within the BGTL refinery for all case studies are shown in Table 6. Biomass may be converted to synthesis gas via gasification that inputs either a solid fuel source (i.e., biomass) or a solid/vapor fuel source (i.e., biomass and recycle gases). Note that the recycle gases could include CO2 that may be consumed via the reverse water gas shift reaction to create additional CO within the biomass gasifier. For all 48 case studies, the gasifier operated using only biomass as the input feedstock, and no recycle gases were selected. This is a direct consequence of the H2/CO ratio that is needed for syngas conversion and the limited amount of H2 within the gasifier. Though CO2 recycle to the gasifier may be a viable alternative if additional H2 was input from a non-carbon based source (e.g., electrolysis), this option is economically unfavorable when electricity prices are high. Alternatively, all case studies employed a CO2 recycle stream to the natural gas conversion unit. The high hydrogen to carbon ratio of natural gas will produce a syngas stream that contains a H2/CO ratio that is usually higher than that needed for synthesis gas conversion. Depending on the relative ratio of natural gas syngas to biomass syngas, the two syngas streams may be combined to have an appropriate H2/CO ratio. For all of the case studies analyzed in this paper, the amount of natural gas syngas was high enough to cause the resulting H2/CO ratio to be above the target of approximately 2. Thus, CO2 recycle to the natural gas conversion unit helped to decrease this ratio to a feasible operational target while simultaneously decreasing process emissions. Natural gas could be converted to syngas via steam reforming (SMR), autothermal reforming, partial oxidation to methanol, or oxidative coupling to olefins. Three possible temperature options were used for the steam reformer (700, 800, 900 °C), the autothermal reformer (800, 900, 1000 °C), and the reverse water gas shift unit (400, 500, 600 °C). For all of the case studies, the natural gas conversion unit was selected to be autothermal reforming. This finding is in contrast to the results for a pure natural gas to liquids (GTL) system that identifies steam reforming as the economically superior technology at lower capacities.33 The key reason for the pervasive use of the autothermal reformer is the presense of the oxygen-blown biomass gasifier in all case studies. Though the capital cost of an autothermal reformer is about half that of a steam reformer, the high capital cost of an air separation unit makes the steam reformer less capital intensive at lower capacities for GTL systems. However, the oxygen requirement for the biomass

Table 5. Cost Parameters (2011 Dollars) for the BGTL Refinery item forest residues perennial residues butanes electricity freshwater

cost $70/dry metric ton $100/dry metric ton $1.84/gal $0.07/kWh $0.50/metric ton

item

cost

agricultural residues natural gas

$120/dry metric ton $5.4/TSCFa

propanes CO2 TS&Mb

$1.78/gal $5/metric ton

a

TSCF, thousand standard cubic feet. bTS&M, transportation, storage, and monitoring.

freshwater, and butanes) include all costs associated with delivery to the plant gate. The products (i.e., electricity and propane) are assumed to be sold from the plant gate and do not include the costs expected for transport to the end consumer. The cost of CO2 capture and compression will be included in the investment cost of the BGTL refinery, while the cost for 3391

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3392

biomass conv biomass temp NG conv NG conv temp WGS/RGS temp min wax FT nom wax FT FT upgrad MTG use MTOD use CO2SEQ use GT use

biomass conv biomass temp NG conv NG conv temp WGS/RGS temp min wax FT nom wax FT FT upgrad MTG use MTOD use CO2SEQ use GT use

biomass conv biomass temp NG conv NG conv temp WGS/RGS temp min wax FT nom wax FT FT upgrad MTG use MTOD use CO2SEQ use GT use

S 1100 ATR 1000 − − − − Y − Y −

S 1100 ATR 1000 − − − − Y − Y −

U-5

U-1

U-5

S 900 ATR 1000 − − − − Y − − −

U-1

S 900 ATR 1000 − − − − Y − − −

S 1100 ATR 1000 − − − − Y − Y −

U-5

S 1100 ATR 1000 − − − − Y − Y −

U-1

S 1000 ATR 1000 − − − − Y − Y −

U-10

S 900 ATR 1000 − − − − Y − − −

U-10

S 1000 ATR 1000 − − − − Y − Y −

U-10

S 1000 ATR 1000 − − − − Y − Y −

U-50

S 900 ATR 1000 − − − − Y − − −

U-50

S 1000 ATR 1000 − − − − Y − Y −

U-50

S 1100 ATR 1000 − − − − − Y Y −

D-1

S 900 ATR 1000 − − − − − Y − −

D-1

S 1100 ATR 1000 − − − − − Y Y −

D-1

S 1100 ATR 1000 − − − − − Y Y −

D-5

S 900 ATR 1000 − − − − − Y − −

D-5

S 1100 ATR 1000 − − − − − Y Y −

D-5

S 1000 ATR 1000 − − − − − Y Y −

D-10

S 900 ATR 1000 − − − − − Y − −

D-10

S 1000 ATR 1000 − − − − − Y Y −

D-10

K-1

S 1100 ATR 1000 − − Co-LTFT fract − − Y − Forest Residues

K-1

Perennial Crops

S 1000 ATR 1000 − − − − − Y Y −

D-50

S 1100 ATR 1000 − − Co-LTFT fract − − Y −

K-1

S S 900 900 ATR ATR 1000 1000 − − − − − Co-LTFT − fract − − Y − − − − − Agricultural Residues

D-50

S 1000 ATR 1000 − − − − − Y Y −

D-50

Table 6. Topological Information for the Optimal Solutions for the Forty-Eight Case Studiesa K-5

S 1100 ATR 1000 − − Co-LTFT fract − − Y −

K-5

S 900 ATR 1000 − − Co-LTFT fract − − − −

K-5

S 1100 ATR 1000 − − Co-LTFT fract − − Y −

K-10

S 1000 ATR 1000 − − Co-LTFT fract − − Y −

K-10

S 900 ATR 1000 − − Co-LTFT fract − − − −

K-10

S 1000 ATR 1000 − − Co-LTFT fract − − Y −

K-50

S 1000 ATR 1000 − − Co-LTFT fract − − Y −

K-50

S 900 ATR 1000 − − Co-LTFT fract − − − −

K-50

S 1000 ATR 1000 − − Co-LTFT fract − − Y −

R-1

S 1100 ATR 1000 − − − − Y Y Y −

R-1

S 900 ATR 1000 − − − − Y Y − −

R-1

S 1100 ATR 1000 − − − − Y Y Y −

R-5

S 1100 ATR 1000 − − Co-LTFT ZSM-5 Y − Y −

R-5

S 900 ATR 1000 − − Co-LTFT ZSM-5 Y − − −

R-5

S 1100 ATR 1000 − − Co-LTFT ZSM-5 Y − Y −

R-10

S 1000 ATR 1000 − − Co-LTFT ZSM-5 Y − Y −

R-10

S 900 ATR 1000 − − Co-LTFT ZSM-5 Y − − −

R-10

S 1000 ATR 1000 − − Co-LTFT ZSM-5 Y − Y −

R-50

S 1000 ATR 1000 − − Co-LTFT ZSM-5 Y − Y −

R-50

S 900 ATR 1000 − − Co-LTFT ZSM-5 Y − − −

R-50

S 1000 ATR 1000 − − Co-LTFT ZSM-5 Y − Y −

Industrial & Engineering Chemistry Research Article

dx.doi.org/10.1021/ie3024643 | Ind. Eng. Chem. Res. 2013, 52, 3381−3406

Article

gasifier enforces the use of an air separation unit in the BGTL refinery. Thus, the incremental capital cost of oxygen production for the autothermal reformer is substantially less than that of a GTL refinery. The temperature of the autothermal reformer is consistently selected to be 1000 °C, which helps to increase both (i) methane conversion in the unit via steam reforming and (ii) CO2 conversion via the reverse water gas shift reaction. Selection of the high-temperature units suggests that the higher species conversions within the reformer due to higher temperature outweights the increased operating costs with a higher temperature. The biomass gasifier could operate at a temperature of 900, 1000, or 1100 °C. For the perennial crop and agricultural residues, the gasifier temperature was selected to be 1100 °C at lower capacities (1 and 5 kBD) while a temperature of 1000 °C was selected for 10 and 50 kBD. The hardwood residues consistently used a gasifier temperature of 900 °C for all case studies. The selection of gasifier temperature for each case study illustrates an important trade-off in utility cost and capital cost both for the biomass gasifier and for other process units. The hardwood residue case studies required a higher amount of energy for biomass drying due to the 45% moisture content assumed for that feedstock. The availability of waste heat for utility generation was therefore not as large as for the two other biomass feedstocks, and the use of a lower temperature gasifier unit could help to partially offset the oxidizer requirement. None of the case studies utilized a dedicated reverse water gas shift unit for CO2 consumption. The equilibrium constant for the water gas shift reaction at the expected operating temperatures of the dedicated unit make for less favorable conditions than the operating temperatures of the reformers or the biomass gasifier. Alternatively, CO2 consumption did occur in the methanol synthesis units. The reverse water gas shift reaction was able to occur at these lower temperatures due to the consumption of CO for the synthesis reactions. This decrease of CO provides the key driver for the consumption of CO2 that is otherwise unavailable in a dedicated reverse water gas shift unit. The selection of hydrocarbon conversion and hydrocarbon upgrading units was largely directed by the type of fuels desired from the BGTL refinery. The 12 case studies that allowed for an unrestricted liquid product composition all selected methanol synthesis and methanol-to-gasoline (MTG) as the optimal technology. This reflects the expected reduction in capital costs associated with hydrocarbon production via methanol synthesis versus FT synthesis that come from the reduced capital cost of methanol synthesis and MTG. Note that gasoline can be produced from FT synthesis and subsequent conversion of the hydrocarbons to gasoline via a ZSM-5 catalyst, but this process requires a higher capital investment over methanol synthesis. Both the MTG and the FT/ZSM-5 processes will produce a significant amount of byproduct liquefied petroleum gas (LPG; 9 vol %). The 12 case studies that maximize the diesel production utilized the methanol-to-olefins (MTO) and the Mobil olefinsto-gasoline/distillate (MOGD) processes to produce a high quality diesel, while the 12 case studies that maximize kerosene will use a cobalt-based low-temperature FT synthesis followed by standard fractionation of the hydrocarbon species. Alternatively, the 12 case studies that produce liquid fuels in the ratios consistent with the demand in the United States show a significant topological trade-off at different capacity levels. That is, at the 1 kBD capacity, methanol synthesis and

a Biomass conversion (biomass conv) is gasification with a solid (S) or solid/vapor (S/V) fueled system. The natural gas conversion technology (NG conv) is either steam reforming (SMR), autothermal reforming (ATR), partial oxidation to methanol (PO), or oxidative coupling to olefins (OC). The temperature (temp; °C) of the biomass/natural gas conversion technology is selected along with the operating temperature of the reverse water gas shift unit (RGS), if utilized. The presence of a CO2 sequestration system (CO2SEQ) or a gas turbine (GT) is noted using “yes” (Y) or “no” (−). The minimum wax and maximum wax Fischer−Tropsch units are designated as either cobalt-based or iron-based units. The iron-based units will either facilitate the forward (fWGS) or reverse water gas shift (rWGS) reaction. The FT vapor effluent will be upgraded (FT upgrad) using fractionation into distillate and naphtha (fract) or ZSM-5 catalytic conversion. The use of methanol-to-gasoline (MTG) and methanol-to-olefins/olefins-to-gasoline-and-diesel (MTO/MOGD) is noted using “yes” (Y) or “no” (−).

Table 6. continued

Industrial & Engineering Chemistry Research

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dx.doi.org/10.1021/ie3024643 | Ind. Eng. Chem. Res. 2013, 52, 3381−3406

Industrial & Engineering Chemistry Research

Article

Figure 6. Process flow diagram for case study K-10.

and turbines are not shown. The PFD shows the biomass conversion through the gasifier and natural gas conversion through the autothermal reformer with oxygen provided by the air separation unit. The biomass gasifier effluent is cleaned to remove residual tar and NH3 (not shown), and both the biomass and natural gas syngas are directed to a precombustion CO2 capture unit. A dual-capture technology is selected for all 48 case studies to decrease the cost of the sulfur recovery system (Claus process) by avoiding the buildup of CO2 in the sulfur conversion loop. This upstream series of conversion units is selected for all of the 48 case studies, which implies that a majority of the CO2 handling is implemented upstream of the hydrocarbon conversion units. The cobalt-based FT units shown in Figure 6 will not produce CO2 as a byproduct, and the input syngas has an H2/CO ratio of 2. The case studies that utilize methanol synthesis will also input a CO2 free syngas that has an H2/CO ratio of 2. The low composition of CO2 effectively means that the CO2 and H2O concentrations in the effluent of the methanol synthesis unit are very low (