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Computers and Chemical Engineering 47 (2012) 29– 56

Contents lists available at SciVerse ScienceDirect

Computers and Chemical Engineering

jo u rn al hom epa ge : www.elsev ier .com/ locate /compchemeng

rocess synthesis of hybrid coal, biomass, and natural gas to liquids viaischer–Tropsch synthesis, ZSM-5 catalytic conversion, methanol synthesis,ethanol-to-gasoline, and methanol-to-olefins/distillate technologies

ichard C. Baliban, Josephine A. Elia, Vern Weekman, Christodoulos A. Floudas ∗

epartment of Chemical and Biological Engineering, Princeton University, Princeton, NJ 08544, USA

r t i c l e i n f o

rticle history:eceived 2 April 2012eceived in revised form 24 June 2012ccepted 25 June 2012vailable online 3 July 2012

eywords:

a b s t r a c t

Several technologies for synthesis gas (syngas) refining are introduced into a thermochemical basedsuperstructure that will convert biomass, coal, and natural gas to liquid transportation fuels usingFischer–Tropsch (FT) synthesis or methanol synthesis. The FT effluent can be (i) refined into gasoline,diesel, and kerosene or (ii) catalytically converted to gasoline and distillate over a ZSM-5 zeolite. Methanolcan be converted using ZSM-5 (i) directly to gasoline or to (ii) distillate via olefin intermediates. A mixed-integer nonlinear optimization model that includes simultaneous heat, power, and water integration

rocess synthesis with heat, power, andater integrationybrid energy systemsixed-integer nonlinear optimization

ischer–Tropschethanol to gasoline

is solved to global optimality to determine the process topologies that will produce the liquid fuels atthe lowest cost. Twenty-four case studies consisting of different (a) liquid fuel combinations, (b) refin-ery capacities, and (c) superstructure possibilities are analyzed to identify important process topologicaldifferences and their effect on the overall system cost, the process material/energy balances, and thewell-to-wheel greenhouse gas emissions.

© 2012 Elsevier Ltd. All rights reserved.
ethanol to olefins and distillate
. Introduction

The transportation sector relies heavily on petroleum as a pri-ary energy source, and currently faces challenges over high crude

il prices, volatility of the global oil market, and greenhouse gasmissions. These issues may be further compounded if the con-ern of domestic peak oil production is to be realized within theoming decades (Nashawi, Malallah, & Al-Bisharah, 2010; Zittel &chindler, 2007). As a result, a substantial effort has been taken inhe United States to focus on greater energy independence throughhe utilization of a more diverse array of primary energy sources. Ofarticular interest are liquid fuels that can be derived from domes-ic carbon-based feedstocks which can directly replace a proportionf the fuels that are produced from crude imports from undesirabler unstable governments (Weekman, 2010). The United States hashree major carbon-based feedstocks, namely coal, biomass, andatural gas, which can be converted into liquid transportation fuelshrough a variety of means. A recent review has highlighted therocess design alternatives that can produce gasoline, diesel, and
erosene using any one or a combination of the three feedstocksFloudas, Elia, & Baliban, 2012).
∗ Corresponding author. Tel.: +1 609 258 4595; fax: +1 609 258 0211.E-mail address: [email protected] (C.A. Floudas).

098-1354/$ – see front matter © 2012 Elsevier Ltd. All rights reserved.ttp://dx.doi.org/10.1016/j.compchemeng.2012.06.032

Each of the three feedstocks has significant advantagesand disadvantages when used to generate liquid transportationfuels. Coal is beneficial because the delivered cost is generallycheaper ($2.0–2.5/MM Btu) (Energy Information Administration,2011) than natural gas ($4.8–5.8/MM Btu) (Energy InformationAdministration, 2011) or biomass ($4.0–9.0/MM Btu) (Kreutz,Larson, Liu, & Williams, 2008; Larson, Jin, & Celik, 2009; NationalAcademy of Sciences, 2009). The prices of all feedstocks are con-verted to 2009$ using the GDP deflator index (US GovernmentPrinting Office, 2009). However, the high carbon content of coalmay require that a significant portion of the feedstock carbon isconverted to CO2 where it can either be vented, sequestered, orconverted back to CO using a non-carbon based source of hydrogen(Agrawal, Singh, Ribeiro, & Delgass, 2007; Baliban, Elia, & Floudas,2010, 2011, 2012; Elia, Baliban, & Floudas, 2010, 2012; Elia, Baliban,Xiao, & Floudas, 2011). Natural gas provides a high hydrogen tocarbon ratio that can help to increase the conversion rates of feed-stock carbon to final liquid fuels. Recent prospects for shale gasproduction have helped reduce the delivered cost of natural gasand made this feedstock a more attractive choice for liquid fuelsproduction (Energy Information Administration, 2011). Biomass ishighly beneficial since it is a renewable energy source that can
absorb atmospheric CO2 during photosynthesis (Lynd et al., 2009;National Academy of Sciences, 2009; Water Science & TechnologyBoard, 2008). Though corn-based ethanol and soybean-based dieselcomprise a majority of the biofuels manufactured today, their use
dx.doi.org/10.1016/j.compchemeng.2012.06.032

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3 Chem

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0 R.C. Baliban et al. / Computers and

or fuel production has led to concerns regarding the impact onhe price and availability of these feedstocks as sources of foodLynd et al., 2009). Lignocellulosic plant sources (e.g., corn stoverr forest residue) are expected to be a more considerable sourcef biofuels in the future, though an increase in crop productionill be required to generate an appropriate amount of sustainable

esidue for fuels production (de Fraiture, Giordano, & Liao, 2008;epartment of Energy, 2005; National Research Council, 2008).

Current alternative energy processes in the literature focus onsing a single feedstock or a hybrid combination of feedstocks toombine the benefits of each feed (Floudas et al., 2012). Generally,hese designs have focused on generation of synthesis gas (syn-as) from the raw materials which is subsequently converted tohe final liquid fuels via the Fischer–Tropsch (FT) process. Exam-les of hybrid systems are coal and biomass to liquids (Agrawalt al., 2007; Chen, Adams, & Barton, 2011a, 2011b; Kreutz et al.,008), coal and natural gas to liquids (Adams & Barton, 2010; Caot al., 2008; Sudiro & Bertucco, 2007, 2009), biomass and natu-al gas to liquids, and coal, biomass, and natural gas to liquidsBaliban, Elia, & Floudas, 2012; Baliban, Elia, Misener, & Floudas,012; Baliban et al., 2010, 2011; Elia et al., 2010). Polygenerationystems have also been designed which are able to co-producelectricity and other liquid fuels along with gasoline, diesel, anderosene (Adams & Barton, 2011; Baliban, Elia, & Floudas, 2012;aliban, Elia, Misener, et al., 2012; Baliban et al., 2011; Chent al., 2011a, 2011b; Chiesa, Consonni, Kreutz, & Williams, 2005;reutz, Williams, Consonni, & Chiesa, 2005; Liu, Pistikopoulos, & Li,009, 2010a, 2010b; Sudiro, Bertucco, Ruggeri, & Fontana, 2008).

n addition to the FT process, liquid fuels can be generated fromyngas by first converting the syngas to methanol and then con-erting the methanol to liquid hydrocarbons (Keil, 1999; Tabak &urchak, 1990). The methanol to gasoline (MTG) (Mobil Research &evelopment Corporation, 1978), methanol to olefins (MTO), andobil olefins to gasoline/distillate (MOGD) (Tabak, Krambeck, &arwood, 1986; Tabak & Yurchak, 1990) processes developed byobil in the 1970s and 1980s utilize a zeolite catalyst to con-

ert the methanol to gasoline and distillate range hydrocarbons.o date, the major process designs involving methanol conversionre based on single feedstock systems including coal to liquids andiomass to liquids (Mobil Research & Development Corporation,978; National Renewable Energy Laboratory, 2011). No hybrideedstock design for methanol synthesis and conversion to liquiduels has been considered.

The previous studies typically focus on process designs wherehe topology of the process is fixed. A process simulation is thenonducted to determine the heat and mass balances for the processnd an economic analysis is performed to determine the viabilityf the plant (Adams & Barton, 2010, 2011; Baliban et al., 2010;echtel Corp, 2003; Cao et al., 2008; Chen et al., 2011a, 2011b;hiesa et al., 2005; Elia et al., 2010; Kreutz et al., 2005, 2008; Larson

Jin, 1999; Liu et al., 2009, 2010a, 2010b; Sudiro & Bertucco, 2007,009; Vliet, Faaij, & Turkenburg, 2009). Recently, optimization-ased process synthesis strategies have been developed where auperstructure of process topologies is postulated and the designhich produces the liquid fuels at the lowest cost or highest profit

s selected (Baliban, Elia, & Floudas, 2012; Baliban, Elia, Misener,t al., 2012; Baliban et al., 2011; Martin & Grossmann, 2011). A com-rehensive process synthesis strategy was proposed which uses aybrid feedstock of coal, biomass, and natural gas to produce theesired liquid fuels (CBGTL) (Baliban, Elia, & Floudas, 2012; Balibant al., 2011), though the methodology could be tailored to ana-yze any one feedstock or combination of feedstocks. A rigorous
lobal optimization framework (Baliban, Elia, Misener, et al., 2012)as also introduced to theoretically guarantee that the cost (orrofit) achieved by the optimal design is within a small percent-ge of the best value possible. The process synthesis strategy is
ical Engineering 47 (2012) 29– 56

capable of directly examining the technoeconomic, environmen-tal, and process topological trade-offs between each design andoutput the solution with the best economic value. The process syn-thesis framework was enhanced by including a simultaneous heatand power integration (Baliban et al., 2011) using an optimization-based heat-integration approach (Duran & Grossmann, 1986) anda series of heat engines that can convert waste heat into electric-ity (Baliban et al., 2010, 2011; Elia et al., 2010). Additionally, theprocess synthesis model integrated a comprehensive wastewatertreatment network (Baliban, Elia, & Floudas, 2012) that utilized asuperstructure approach (Ahmetovic & Grossmann, 2010a, 2010b;Grossmann & Martín, 2010; Karuppiah & Grossmann, 2006) todetermine the appropriate topology and operating conditions ofprocess units that are needed to minimize wastewater contami-nants and freshwater intake.

This manuscript introduces several distinct methods for conver-sion of syngas to liquid fuels into the CBGTL process superstructureand investigates the tradeoffs that arise from these methods. Theprevious superstructure converted the syngas into a raw FT hydro-carbon product using one of four FT units operating with eithera cobalt or iron catalyst and at high or low temperature (Baliban,Elia, & Floudas, 2012; Baliban et al., 2011). The effluent was subse-quently fractionated and upgraded using a series of hydrotreatingunits, a wax hydrocracker, two isomerization units, a naphthareformer, an alkylation unit, and a gas separation plant (i.e., deeth-anizer). This study introduces two iron-based FT units that utilizethe forward water–gas-shift reaction to produce the raw hydrocar-bons using an input H2/CO ratio that is less than the typical 2/1ratio needed for FT synthesis. Catalytic conversion of the FT vaporeffluent over a ZSM-5 catalyst is considered as an alternative forproducing gasoline range hydrocarbons from the raw FT effluent.Methanol synthesis and subsequent conversion to liquid hydro-carbons are also introduced into the superstructure. The methanolmay be catalytically converted using a ZSM-5 zeolite to (i) gasolinerange hydrocarbons or (ii) to distillate (i.e., diesel and kerosene)via an intermediate coversion to olefins. The mathematical model-ing and cost functions needed to incorporate the above alternativesinto the superstructure are outlined in detail. The complete processsynthesis optimization model is then tested on a total of 24 casestudies which consist of two liquid product combinations, threeplant capacities, and four plant superstructures. Using low-volatilebituminous coal (Illinois #6) and perennial biomass (switchgrass),important topological differences between the case studies are dis-cussed and the results of each component of the process synthesisframework are illustrated.

2. CBGTL mathematical model for process synthesis withsimultaneous heat, power, and water integration

This section will discuss the enhancements to the previousmathematical model for process synthesis and simultaneous heat,power, and water integration that will incorporate a wide vari-ety of designs for syngas conversion and hydrocarbon upgrading.Modeling of these enhancements will be described in detail in thefollowing section and the complete mathematical model is listed inAppendix A. The nomenclature used in the mathematical descrip-tion below is outlined in Table 1. Note that this table represents asubset of the comprehensive list of symbols that are needed for thefull mathematical model. The full list of symbols and mathematicalmodel are included for reference in Appendix A.

2.1. Conceptual design

The syngas conversion and hydrocarbon upgrading units pro-posed in this paper is based on an extension of the CBGTL

R.C. Baliban et al. / Computers and Chemical Engineering 47 (2012) 29– 56 31

HTFT

Cob alt LT Min-

Wax FT

Clean Gas

(from AGR )

LTFT

Cob alt HT Nom -

Wax FT

LTFTRG S

Iron HT rWGS

Nom -Wax FT

Hydrog en

(from XH2-HTFT)

PSA Offgas

(from SPPSA)

CO2

(from XCO2-HTFT)

Reformed Gases

(from SPATR)PSA Offgas

(from SPPSA) HTFTRG S

Iron LT rWGS

Min-Wa x FT

Hydrog en

(from XH2-HTFTRGS)

Reformed Gases

(from SPATR)

MTFTWGS-N

Iron MT fWGS

Nom -Wax FT

Hydrog en

(from XH2-MTFTN)

Steam

(from SPSTM)

FT Hydrocarbons

(to SPFTH)Wax

(to WHC)

MTFTWGS-M

Iron -MT fWGS

Min-Wa x FT

Hydrog en

(from XH2-MTFTM)

Steam

(from SPSTM)

FT Hydrocarbons

(to SPFTH)

Wax

(to WHC)

FT Hydrocarbons

(to SPFTH)

FT Hydrocarbons

(to SPFTH)

Hydrog en

(from XH2-LT FT)

Hydrog en

(from XH2-HTFT)

SPCG

SPFTM

SPFTN

SPFTHT

SPFTL T

CO2

(from XCO2-LT FT)

CO2 Lean

Light Gases

(from CO2SEP)

Clean Gas

(to MEOH S)

Fig. 1. Fischer–Tropsch (FT) hydrocarbon production flowsheet. Each of the six FT units has a distinct set of operating conditions including catalyst type (cobalt or iron),temperature (low – 240 ◦C, medium – 267 ◦C, and high – 320 ◦C), and water–gas-shift reaction extent (forward, reverse, or none). Each unit is designed to produce either aminimal or nominal amount of wax (shown as a dashed line). The mathematical model will select at most two types of the six FT units to operate in a final process topology.

WSOS

Water Solub le

Oxygenates Sep.

Water Lean FT

Hydrocarbons

(to HRC )

VLWS

Vapo r / Liqu id /

Water Separator

Wastewater

(to MXSS)

VPOS

Vapo r Phase

Oxygenates Sep.

OxygenatesOxygenates

FT Hydrocarbons

FT-ZSM5

ZSM-5 UnitDistill ate

(to DHT)

Sour Wa ter

(to MXPUWW)

ZSM5F

ZSM-5 Produ ct

Frac tion ation

Raw ZSM-5

HC Produ ct

(to SPFTZSM)

SPFTH

MXFTWW

MXHRC

Raw

Produ ct

Fig. 2. First Fischer–Tropsch (FT) hydrocarbon upgrading flowsheet. The FT effluent may be passed through a series of stripper and flash units to separate the oxygenatesand aqueous phase from the hydrocarbons. Alternatively, the effluent may be passed over a ZSM-5 catalytic reactor to convert most of the hydrocarbons into gasoline rangespecies. The raw ZSM-5 product is then fractionated to remove any distillate or sour water from the gasoline product.

32 R.C. Baliban et al. / Computers and Chemical Engineering 47 (2012) 29– 56

7

WHC

Wax

Hydrocrac ker

Sour Wa ter

(to MXPUWW)

NRF

Naph tha

Reformer

Water Lean FT

Hydrocarbons

(from MXHRC) HRC

Hydroca rbon

Rec overy System NHT

Naph tha

Hydrotrea ter

DHT

Distill ate

Hydrotrea ter

Light Off gas

Naph tha

Hydrog en

(from SPH2)

C56I

C5/C6

Isomerizer

C4I

C4 Isomerizer

C345A

C3/C4/C5

Alkyli zer

SGP

Saturated Gas

Plant

C4 Gases

C4 Gases

Isomerized C4

GasesLight

Off gas

C3/C4/C5

Gases

GBL

Gasoli ne

Blend er

DBL

Diesel

Blend er

Off gas (fr om other unit s)

Sour Wa ter

(to MXPUWW )

Hydrog en

(from SPH2)

Off gas

(to SGP)

Off gas

(to SGP)

Diesel

Sour Wa ter

(to MXPUWW )

Off gas

(to SGP)

Sour Wa ter

(to MXPUWW )

Hydrog en

(from SPH2)

Hydrog en

(from SPH2)

Distill ate

Kerosene

Hydrog en

(from SPH2)

Off gas

(to SGP)Off gas

(to SGP)

Isomerate

Reformate

C5/C6

Gases

Alky late

Trea ted

Naph tha

C5/C6 Gases

Crac ked

Naph tha

Diesel

Light Off gas

OUTLPG

Outpu t LPG

OUTGAS

Outpu t

Gasoline

OUTDIE

Outpu t

Diesel

INBUT

Inpu t

Butane

SPLPG

Light Gases

(to SPLG)MXSGP

Wax

(from MXWAX)

KHT

Kerosene

Hydrotrea ter

Hydrog en

(from SPH2)

Off gas

(to SGP)

OUTKER

Outpu t

Kerosene

Distill ate

Kerosene

Fig. 3. Second Fischer–Tropsch (FT) hydrocarbon upgrading flowsheet. The water lean FT effluent is fractionated and passed through a series of treatment units to recoverthe gasoline, diesel, and kerosene products along with some LPG byproduct. Light gases (i.e., unreacted syngas and C1–C2 hydrocarbons) are collected and recycled back tothe process.

Clean Gas

(from SPCG)

MEOH S

Methano l

Synthesis

H2

(from SPH2)

Raw Methano l

Produ ct

Purified

Methanol

MTG

Methano l to

Gasoline

Raw MTOD

HC Produ ct

(to SPMTODHC)

MEDEG

Methano l

Degass er

Offgas

(to SPLG)

MTO

Methano l to Olefins

MTODF

Hydroca rbon

Frac tion ationRaw

Produ ct

OGD

Olefins to Gasoli ne

and Distil late

Ker osene

(to KH T)

Distill ate

(to DHT)Raw

Olefins

Raw MTG HC

Produ ct

(to SPMTGHC)

MTO-F

Olefin Fraction ation

Olefins

Light Gases

(to SPLG)

OUTGAS

Outpu t Gasoline

Wastewater

(to MXPUWW)

Wastewater

(to MXPUWW)

SPMEOH

MEOH-F

Methano l Flash

Offgas

(to SPLG)

Fig. 4. Methanol synthesis and conversion flowsheet. Clean syngas is initially converted to methanol and then split to either the methanol to gasoline (MTG) or methanol toolefins (MTO) processes. The two processes utilize a ZSM-5 zeolite to convert the methanol to either gasoline range hydrocarbons (MTG) or olefins which are subsequentlyoligomerized to gasoline and distillate range hydrocarbons (MOGD). The distillate is hydrotreated to form diesel or kerosene which the gasoline range hydrocarbons are sentto an LPG-gasoline separation system.


HCKO 2

Mixed HC

Cond ensation

Kno ckou t 2

Light gases

OUTGAS

Outpu t Gasoline

OUTLPG

Outpu t LPG

DEET H

Dee thanizer

ABS-COL

Absorber Column

Reflux

Light gases

STA-COL

Stabili zer Column

SP-COL

Split ter Column

ALK-UN

HF Alkylation

Unit

LPG-ALK

LPG / Alkylate

Split ter

Crud e

HC

C3+ HC

Light HC

gas

Lea n oil

rec ycle

C3/C4 gases

LPG /

Alky late

INBUT

Inpu t Butane

CO2SEP

CO2 Separation Unit

Recovere d CO2

(to MXCO2C)

CO2 Lean

Light ga ses

(to SPCG)

Raw ZSM-5 HC

Produ ct

(from ZSM5F)

Raw MTG

HC Produ ct

(from MTG)

HCKO 1

Mixed HC

Cond ensation

Kno ckou t 1

CO2 Rich

Light ga ses

(to SPLG)

Raw MTOD

HC Produ ct

(from MTODF)

SPFTZSM

SPMTGHC

SPMTODHC

Fig. 5. LPG-gasoline product separation flowsheet. The raw HC products from the FT-ZSM5 unit, the MTG unit, or the MOGD process are passed through a series of separationu ack tor

r(tcaPBetMliati

TM

nits to recover a gasoline product and an LPG byproduct. Light gases are recycled beaction with H2 via the reverse water–gas-shift reaction.

efinery superstructure detailed by Baliban, Elia, and Floudas2012), Baliban et al. (2011). All relevant thermodynamic informa-ion (i.e., chemical equilibrium constants, vapor-liquid equilibriumonstants, specific enthalpies, and heat capacities) for the unitsnd streams in the refinery have been extracted from Aspenlus v7.3 using the Peng–Robinson equation of state with theoston–Mathias alpha function. The flowsheets depicting thextensions of the superstructure are shown in Figs. 1–5 of thisext and the complete superstructure is included as Supplementary

aterial. In the figures, fixed process units are represented in aight blue color, variable process units in dark blue, splitter unitsn gold, and mixer units in green. Fixed process streams will have

gray color while the variable process streams are in blue. Notehat some units (e.g., compressors, pumps, heat exchangers) are notncluded in the figures for clarity, though these units are thoroughly

able 1athematical model nomenclature.

Symbol Definition

Indicess Species indexu Process unit index

Sets(u, u′) Stream from unit u to unit u′

(u, u′ , s) Species s within stream (u, u′)u ∈ UIr

FTSet of all iron-based FT units

ParametersKWGS

u Water–gas-shift equilibrium constant for unit uKMSN

u Methanol synthesis equilibrium constant for unit u

VariablesNS

u,u′,s Molar flow of species s from unit u to unit u′

xSu,u′,s Molar concentration of species s from unit u to unit u′

the refinery and CO2 recovery may be utilized in preparation for sequestration or

modeled in the CBGTL refinery, as noted in previous studies(Baliban, Elia, & Floudas, 2012; Baliban, Elia, Misener, et al., 2012;Baliban et al., 2011).

The CBGTL superstructure is designed to co-feed biomass, coal,or natural gas to produce gasoline, diesel, and kerosene. Syngasis generated via gasification from biomass (Supp. Fig. S1) or coal(Supp. Fig. S2) or auto-thermal reforming of natural gas (Supp. Fig.S10). Co-feeding of the coal, biomass, or natural gas in a singlegasifier unit was not considered in this study due to the lack of(i) technical maturity of the process design and (ii) cost and oper-ating data for co-fed units. Synergy for co-fed biomass and coalgasification and simultaneous reforming the natural gas using thegasifier quench heat (Adams & Barton, 2011) may be important toreduce the capital cost required for synthesis gas production, andthe authors note that the optimization model is capable of includ-ing the technoeconomic benefit of co-fed gasification if cost andoperational data become available.

The synthesis gas is either (i) converted into hydrocarbon prod-ucts in the Fischer–Tropsch (FT) reactors (Fig. 1; Supp. Fig. S5) or (ii)into methanol via methanol synthesis (Fig. 4; Supp. Fig. S8). The FTwax will be sent to a hydrocracker to produce distillate and naphtha(Supp. Fig. S7) while the FT vapor effluent may be (a) fractionatedand upgraded into gasoline, diesel, or kerosene or (Fig. 3; Supp.Fig. S7) (b) catalytically converted to gasoline via a ZSM-5 zeolite(Fig. 2; Supp. Fig. S6). The methanol may be either (a) catalyticallyconverted to gasoline via the ZSM-5 catalyst (Figs. 4–5; Supp. Figs.S8–S9) or (b) catalytically converted to olefins via the ZSM-5 cata-lyst and subsequently fractionated to distillate and gasoline (Fig. 4;Supp. Fig. S8).

Acid gases including CO2, H2, and NH3 are removed from thesyngas via a Rectisol unit prior to conversion to hydrocarbons ormethanol (Supp. Fig. S3). Incorporation of other acid gas removaltechnologies (e.g., amine absorption, pressure-swing absorption,

3 Chem

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acuum-swing absorption, membrane separation) and their rela-ive capital/operating cost as a function of input flow rate and acidas concentration is the subject of an ongoing study. The sulfur-richases are directed to a Claus recovery process (Supp. Fig. S4) andhe recovered CO2 may be sequestered (Supp. Fig. S3) or reactedith H2 via the reverse water–gas-shift reaction. The CO2 may

e directed to either the gasifiers (Supp. Figs. S1–S2), the reverseater–gas-shift reactor (Supp. Fig. S3), or the iron-based FT units

Supp. Fig. S5). Recovered CO2 is not sent to the cobalt-based FTnits to ensure a maximum molar concentration of 3% and preventoisoning of the catalyst. Hydrogen is produced via pressure-swingbsorption or an electrolyzer unit while oxygen can be provided byhe electrolyzer or a distinct air separation unit (Supp. Fig. S11). Aomplete water treatment network (Supp. Figs. S12–S13) is incor-orated that will treat and recycle wastewater from various processnits, blowdown from the cooling tower, blowdown from the boil-rs, and input freshwater. Clean output of the network includes (i)rocess water to the electrolyzers, (ii) steam to the gasifiers, auto-hermal reactor, and water–gas-shift reactor, and (iii) dischargedastewater to the environment.

The effluent of each reactor in the CBGTL refinery is basedn either (i) known extents of reaction, (ii) thermodynamically-imited equilibrium, or (iii) a specified composition from a literatureource. Reaction system (i) is used in the gasifiers, the tar cracker,nd the combustor units (e.g., fuel combustor, gas turbine, Clausombustor) and the extents of reaction are based on known infor-ation from literature (gasifiers/cracker) or from the operating

onditions of the unit (i.e., complete combustion using a stoi-hiometric excess of oxygen). Reaction system (ii) is used forhe water–gas-shift reaction (i.e., gasifiers, WGS reactor, FT units,

ethanol synthesis, auto-thermal reactor), methanol synthesis,nd steam reforming in the auto-thermal reactor. Reaction systemiii) is used for the FT units, the ZSM-5 hydrocarbon conversion, the

TG reactor, the MTO reactor, and the MOGD reactor. The effluentomposition of these units is based on known commercial data orilot plant data for the units operating at a specified set of condi-ions (i.e., temperature, pressure, and feed composition). The CBGTLrocess is designed to ensure that the appropriate conditions areet within the reactor to ensure that the effluent composition that

s assumed is valid. Binary decision variables (y) are included withinhe mathematical model to logically define the existence of specificrocess units (Eqs. (A.17)–(A.21)). That is, if y = 0 for a particularnit, then no heat/mass flow will be allowed through the unit andhe unit will effectively be removed from the process topology. If

= 1 for a unit, then the heat/mass flow through the unit will beoverned by the proper operation of the unit.

.2. Fischer–Tropsch units

The four FT units considered in previous studies (Baliban, Elia, Floudas, 2012; Baliban et al., 2011) utilized either a cobalt or ironatalyst and operated at high or low temperature. The two cobalt-ased FT units would not facilitate the water–gas-shift reactionnd therefore required a minimal level of CO2 input to the unitso improve the per–pass conversion of CO. The two iron-based FTnits were assumed to facilitate the reverse water–gas-shift reac-ion and therefore could consume CO2 within the unit using H2 toroduce the CO necessary for the FT reactions. A key synergy ofhe reaction conditions in the latter units was the heat needed forhe reverse water–gas-shift reaction that is provided by the highlyxothermic FT reaction. Though the reverse water–gas-shift reac-ion is typically unfavorable at the lower operating temperatures
f the FT units, the reaction may be indeed facilitated through these of an appropriate amount of input hydrogen.
In this study, the set of possible FT units is expanded to considerron-based systems that will facilitate the forward water–gas-shift


reaction within the units. These FT units will require a lower H2/COratio for the FT reaction because steam in the feed will be shifted toH2 through consumption of CO. These units may be beneficial sincecertain syngas generation units (e.g., coal gasifiers) will producea gas that generally has a H2/CO ratio that is much less than the2/1 requirement for FT synthesis (Baliban et al., 2010; Kreutz et al.,2008). The downside of the new FT units will be the high quantity ofCO2 that is produced as a result of the water–gas-shift reaction. Theframework developed for the CBGTL superstructure will directlyexamine the benefits and consequences for each of the six FT unitsto determine which technology produces a refinery with a superiordesign.

Fig. 1 shows the flowsheet for FT hydrocarbon production withinthe superstructure. Clean gas from the acid gas removal (AGR) unitis mixed with recycle light gases from a CO2 separator (CO2SEP) andsplit (SPCG) to either the low-wax FT section (SPFTM), the nominal-wax FT section (SPFTN), or methanol synthesis (MEOHS). The FTunits will operate at a pressure of 20 bar and within the temper-ature range of 240–320 ◦C. The cobalt-based FT units operate ateither low temperature (LTFT; 240 ◦C) or high temperature (HTFT;320 ◦C) and must have a minimal amount of CO2 in the input stream.Two iron-based FT units will facilitate the reverse water–gas-shift(rWGS) reaction and will operate at low (LTFTRGS; 240 ◦C) and hightemperature (HTFTRGS; 320 ◦C). The other two iron-based FT unitswill use the forward reverse water–gas-shift (fWGS) units, operateat a mid-level temperature (267 ◦C), and produce either minimal(MTFTWGS-M) or nominal (MTFTWGS-N) amounts of wax. Theoperating conditions of the FT units are summarized in Table 2.

Hydrogen may be recycled to any of the FT units to either shiftthe H2/CO ratio or the H2/CO2 ratio to the appropriate level. Steammay alternatively be used as a feed for the two iron-based fWGSFT units to shift the H2/CO ratio. CO2 may be recycled back to theiron-based rWGS FT units to be consumed in the WGS reaction.Similarly, the pressure-swing absorption (PSA) offgas which willbe lean in H2 may be recycled to the iron-based rWGS FT units forconsumption of the CO or CO2. The effluent from the auto-thermalreactor (ATR) will contain a H2/CO ratio that is generally above 2/1,and is therefore favorable as a feedstock for FT synthesis (NationalAcademy of Sciences, 2009). However, the concentration of CO2within the ATR effluent will prevent the stream from being fed tothe cobalt-based units. The two streams exiting the FT units willbe a waxy liquid phase and a vapor phase containing a range ofhydrocarbons. The wax will be directed to a hydrocracker (WHC)while the vapor phase is split (SPFTH) for further processing.

Modeling of the four original FT units has been previouslydescribed (Baliban, Elia, & Floudas, 2012; Baliban et al., 2011) andis included in the Supplementary Material. The effluent from thetwo additional FT units (iron-based FT fWGS) is based off of theslurry phase FT units developed by Mobil Research and Develop-ment Corporation in the 1980s (Mobil Research & DevelopmentCorporation, 1983, 1985). A H2/CO ratio of 2/3 is desired for theinput feed (Mobil Research & Development Corporation, 1983,1985), so a sufficient amount of steam must be added to the feed topromote the forward water–gas-shift reaction. The decompositionof carbon from CO to hydrocarbons and CO2 is outlined in Table42 of the minimal-wax FT report (Mobil Research & DevelopmentCorporation, 1983) and Table VIII-2 of the nominal-wax FT report(Mobil Research & Development Corporation, 1985), and a 90% con-version of CO in the inlet stream is assumed (Mobil Research &Development Corporation, 1983, 1985). The syngas species exitingthe four iron-based FT reactors will be constrained by water–gas-shift equilibrium, as noted in Eq. (1) where (u, u′) is the stream
exiting the FT unit u.
NSu,u′,H2O · NS

u,u′,CO = KWGSu · NS

u,u′,H2· NS

u,u′,CO2∀u ∈ UIr

FT (1)


Table 2Operating conditions for the process units involved in methanol synthesis and conversion to liquid hydrocarbon fuels.

Unit Temperature (◦C) Pressure (bar) Conv.

LT cobalt FT synthesis 240 20 80% of COLT iron FT synthesis 240 20 80% of COMT iron FT synthesis (low wax) 267 20 90% of COMT iron FT synthesis (high wax) 267 20 90% of COHT cobalt FT synthesis 320 20 80% of COHT iron FT synthesis 320 20 80% of COZSM-5 FT upgrading 408 16 100% of hydrocarbonsMethanol synthesis 300 50 30–40% of CO

Fcatu

2

tFsstotaurFfwt

a(iRwsau(VDctCblC(uaTs

hcw

Methanol-to-gasoline 400Methanol-to-olefins 482

Olefins-to-gasoline/distillate 300

The mathematical model will select at most two types ofischer–Tropsch units to operate in the final process design. Thisonstraint is added because two different kinds of FT units will beble to supply a range of hydrocarbon species that is diverse enougho provide a target composition of liquid products without addingnnecessary complexity to the refinery design (de Klerk, 2011).

.3. Fischer–Tropsch product upgrading

The vapor phase effluent from FT synthesis will contain a mix-ure of C1–C30+ hydrocarbons, water, and some oxygenated species.ig. 2 details the process flowsheet used to process this effluenttream. The stream will be split (SPFTH) and can pass through aeries of treatment units designed to cool the stream and knock outhe water and oxygenates for treatment. Initially, the water-solublexygenates are stripped (WSOS) from the stream. The stream ishen passed to a three-phase separator (VLWS) to remove thequeous phase from the residual vapor and any hydrocarbon liq-id. Any oxygenates that are present in the vapor phase may beemoved using an additional separation unit (VSOS). The water leanT hydrocarbons are then sent to a hydrocarbon recovery columnor fractionation and further processing (Fig. 3). The oxygenates andater removed from the stream are mixed (MXFTWW) and sent to

he sour stripper mixer (MXSS) for treatment.The FT hydrocarbons split from SPFTH may also be passed over

ZSM-5 catalytic reactor (FT-ZSM5) operating at 408 ◦C and 16 barMobil Research & Development Corporation, 1983) to be convertednto mostly gasoline range hydrocarbons and some distillate (Mobilesearch & Development Corporation, 1983, 1985). The ZSM-5 unitill be able to convert the oxygenates to additional hydrocarbons,

o no separate processing of the oxygenates will be required for thequeous effluent. The composition of the effluent from the ZSM-5nit is shown in Table 43 of the minimal-wax FT reactor Mobil studyMobil Research & Development Corporation, 1983) and in TableIII-3 of the nominal-wax FT reactor Mobil study (Mobil Research &evelopment Corporation, 1985). For this study, the ZSM-5 effluentomposition is assumed to be equal to the composition outlined inhe minimal-wax FT reactor study (Mobil Research & Developmentorporation, 1983). This is modeled mathematically using an atomalance around the ZSM-5 unit and the effluent composition out-

ined in Table 43 of the Mobil study (Mobil Research & Developmentorporation, 1983). The raw product from FT-ZSM5 is fractionatedZSM5F) to separate the water and distillate from the gasoline prod-ct. The water is mixed with other wastewater knockout (MXPUWW)nd the distillate is hydrotreated (DHT) to form a diesel product.he raw ZSM-5 HC product is sent to the LPG-gasoline separationection for further processing (Fig. 5).

The water lean FT hydrocarbons leaving MXFTWW are sent to aydrocarbon recovery column (HRC), as shown in Fig. 3. The hydro-arbons are split into C3–C5 gases, naphtha, kerosene, distillate,ax, offgas, and wastewater (Baliban et al., 2010; Bechtel, 1998).

12.8 100% of methanol1.1 100% of methanol

50 100% of olefins

The upgrading of each stream will follow a detailed Bechtel design(Bechtel, 1992, 1998) which includes a wax hydrocracker (WHC),a distillate hydrotreater (DHT), a kerosene hydrotreater (KHT), anaphtha hydrotreater (NHT), a naphtha reformer (NRF), a C4 iso-merizer (C4I), a C5/C6 isomerizer (C56I), a C3/C4/C5 alkylation unit(C345A), and a saturated gas plant (SGP).

The kerosene and distillate cuts are hydrotreated in (KHT) and(DHT), respectively, to remove sour water and form the productskerosene and diesel. Any additional distillate or kerosene producedin other sections of the refinery will also be directed to these unitsfor processing. The naphtha cut is sent to a hydrotreater (NHT) toremove sour water and separate C5–C6 gases from the treated naph-tha. The wax cut is sent to a hydrocracker (WHC) where finisheddiesel product is sent to the diesel blender (DBL) along with thediesel product from (DHT). C5–C6 gases from (NHT) and (WHC) aresent to an isomerizer (C56I). Hydrotreated naphtha is sent to thenaphtha reformer (NRF). The C4 isomerizer (C4I) converts in-plantand purchased butane to isobutane, which is fed into the alkyla-tion unit (C345A). Purchased butane is added to the isomerizer suchthat 80 wt% of the total flow entering the unit is composed of n-butane. Isomerized C4 gases are mixed with the C3–C5 gases fromthe (HRC) in (C345A), where the C3–C5 olefins are converted to high-octane gasoline blending stock. The remaining butane is sent backto (C4I), while all light gases are mixed with the offgases from otherunit and sent to the saturated gas plant (SGP). C4 gases from (SGP)are recycled back to the (C4I) and a cut of the C3 gases are sold asbyproduct propane.

2.4. Methanol synthesis and conversion

The clean gas split (SPCG) from the acid gas recovery unit maybe directed to a methanol synthesis unit (MEOHS) for conversionof the syngas to methanol (National Renewable Energy Laboratory,2011). The syngas exiting the acid gas recovery unit is heated up to300 ◦C prior to entering the MEOHS unit. The MEOHS unit operatesat a temperature of 300 ◦C, a pressure of 51 bar, and will assumeequilibrium between the water–gas-shift reaction (Eq. (2)) andthe methanol synthesis reaction (Eq. (3)) in the effluent stream(MEOHS, u) (National Renewable Energy Laboratory, 2011).

NSMEOHS,u,H2O · NS

MEOHS,u,CO = KWGSMEOHS · NS

MEOHS,u,H2· NS

MEOHS,u,CO2

(2)

xSMEOHS,u,CH3OH = KMSN

MEOHS ·(

xSMEOHS,u,H2

)2 · xSMEOHS,u,CO (3)

Note that the equations for water–gas-shift equilibrium (Eqs.
(1) and (2)) utilize molar species flow rates while the methanolsynthesis equilibrium (Eq. (3)) and the steam reforming equilib-rium (Eqs. (A.98)–(A.101)) utilize molar species concentrations. Theconservation of total moles across the water–gas-shift equilibrium

3 Chem

acmutTsrecosieu

umtiotspyortiobi9TtTdas

t1og&w1LyLmao1(lasroa4bgh(


llows for the use of either species molar flow rates or molar con-entrations in the equilibrium reaction without a need for a totalolar flow rate variable. The mathematical model was formulated

sing molar flow rates because the bilinear terms for calculation ofhe concentration variables are not required for all syngas species.he remainder of the chemical equilibrium equations do not con-erve the amount of total moles, so the use of species molar flowates would require a total molar flow rate variable to balance thequation. In this study, it was found to be computationally benefi-ial to use species concentration variables to reduce the presencef trilinear or quadrilinear terms that would arise with the use ofpecies molar flow rates. Note that the equilibrium constants usedn Eqs. (3) and (A.98)–(A.101) have been modified from the valuesxtracted from Aspen to account for the increased pressure of thenits.

The “state-of-technology” conditions for methanol synthesissed in this study will require a CO2 input concentration of 3–8% forethanol synthesis (National Renewable Energy Laboratory, 2011),

hough there could exist a potential synergy from a higher CO2nput concentration (Toyir et al., 2009). However, an increased levelf H2 may also need to be input to the reactor for consumption viahe reverse water–gas-shift reaction. H2 generated via pressure-wing absorption may not be appropriate if the H2-lean offgas isrimarily used as plant fuel. Alternatively, H2 provided by electrol-sis of water with a non-carbon-based form of electricity (e.g., windr solar) will have a high capital cost of electrolyzers coupled with aelatively high cost of renewable-based electricity. This may offsethe reduction in capital that is achieved if a CO2 capture technologys not needed for the synthesis gas. The technoeconomical benefitsf higher levels of CO2 input to the methanol synthesis reactor wille the subject of a future investigation. The raw methanol effluent

s cooled to 35 ◦C and sent to a flash unit (MEOH-F) to remove over5% of the entrained methanol through vapor–liquid equilibrium.he vapor phase is split and mostly recycled (split fraction: 95%) tohe methanol synthesis reactor to increase the yield of methanol.he methanol leaving the MEOH-F unit is degassed (MEDEG) viaistillation to remove any light vapors. The MEDEG unit is oper-ted as a split unit with a steam utility requirement derived throughimulation.

The purified methanol is split (SPMEOH) to either the methanol-o-gasoline (MTG) (Mobil Research & Development Corporation,978; National Renewable Energy Laboratory, 2011) processr to the methanol-to-olefins (MTO) and Mobil olefins-to-asoline/distillate (MOGD) (Keil, 1999; Tabak et al., 1986; Tabak

Krambeck, 1985; Tabak & Yurchak, 1990) processes, both ofhich were developed by Mobil Research and Development in the

970s and 1980s. More recently, the National Renewable Energyaboratory performed a full design, simulation, and economic anal-sis of a biomass-based MTG process (National Renewable Energyaboratory, 2011). The MTG process will catalytically convert theethanol to gasoline range hydrocarbons using a ZSM-5 zeolite and

fluidized bed reactor. The MTG effluent is outlined in Table 3.4.2f the Mobil study (Mobil Research & Development Corporation,978) and in Process Flow Diagram P850-A1402 of the NREL studyNational Renewable Energy Laboratory, 2011). Due to the highevel of component detail provided by NREL for both the MTG unitnd the subsequent gasoline product separation units, the compo-ition of the MTG reactor used in this study is based on the NRELeport. The MTG unit will operate adiabatically at a temperaturef 400 ◦C and 12.8 bar. The methanol feed will be heated to 330 ◦Cnd input to the reactor at 14.5 bar. The MTG effluent will contain4 wt% water and 56 wt% crude hydrocarbons, of which 2 wt% will
e light gas, 19 wt% will be C3–C4 gases, and 19 wt% will be C5+asoline (National Renewable Energy Laboratory, 2011). The crudeydrocarbons are directed to the LPG-gasoline separation sectionFig. 5), from which 82 wt% will be gasoline, 10 wt% will be LPG, and

the balance will be recycle gases. This is modeled mathematically inthe process synthesis model by using an atom balance around theMTG unit and assuming a 100% conversion of the methanol enter-ing the MTG reactor (Mobil Research & Development Corporation,1978; National Renewable Energy Laboratory, 2011).

Any methanol entering the MTO process unit is heated to 400 ◦Cat 1.2 bar. The MTO fluidized bed reactor operates at a tempera-ture of 482 ◦C and a pressure of 1.2 bar (Tabak & Yurchak, 1990).The exothermic heat of reaction within the MTO unit is con-trolled through generation of low-pressure steam. 100% of theinput methanol is converted into olefin effluent containing 1.4 wt%CH4, 6.5 wt% C2–C4 paraffins, 56.4 wt% C2–C4 olefins, and 35.7 wt%C5–C11 gasoline (Tabak & Yurchak, 1990). The MTO unit is modeledmathematically using an atom balance and a typical compositionseen in the literature (Tabak & Yurchak, 1990). The MTO product isfractionated (MTO-F) to separate the light gases, olefins, and gaso-line fractions. The MTO-F unit is a rigorous distillation column thatis designed so that approximately 100% of the C1–C3 paraffins arerecycled back to the refinery, 100% of the C4 paraffins and 100% ofthe olefins are directed to the MOGD unit, 100% of the gasoline iscombined with the remainder of the gasoline generated in the pro-cess, and 100% of the water generated in the MTO unit is sent forwastewater treatment. Note that the MTO-F unit is modeled withinthe process synthesis model as a separator unit with the appropri-ate utilities (i.e., low-pressure steam and cooling water) that areextracted from simulation of the distillation column.

The separated olefins are sent to the MOGD unit where a fixedbed reactor is used to convert the olefins to gasoline and distil-late over a ZSM-5 catalyst. The gasoline/distillate product ratioscan range from 0.12 to >100, and the ratio chosen in this study was0.12 to maximize the production of diesel. The MOGD unit oper-ates at 400 ◦C and 1 bar and will utilize steam generation to removethe exothermic heat of reaction within the unit. The MOGD unit ismodeled with an atom balance and will produce 82% distillate, 15%gasoline, and 3% light gases (Tabak & Yurchak, 1990). The productwill be fractionated (MTODF) to remove diesel and kerosene cutsfrom the gasoline and light gases. The operational ratio of keroseneto total distillate reported in the literature for the MOGD processis about 30%, though this number may be increased by tailoringthe operating conditions within the MTO and MOGD units to yieldthe appropriate range of hydrocarbons. The MTODF unit will bemodeled as a separator unit where 100% of the C11–C13 speciesare directed to the kerosene cut and 100% of the C14+ species aredirected to the diesel cut.

2.5. LPG-gasoline separation

The gasoline range hydrocarbons produced by the FT-ZSM5 unit,the MTG unit, or the MOGD process must be sent to the LPG-gasoline separation flowsheet depicted in Fig. 5. Each hydrocarbonstream is split (SPFTZSM, SPMTGHC, and SPMTODHC, respectively) andsent to a hydrocarbon knockout unit (35 ◦C, 10 bar) for light gasremoval via vapor–liquid equilibrium. The first knock-out unit(HCKO1) will not incorporate additional CO2 separation, so the CO2rich light gases recovered from HCKO1 will be recycled back to theprocess (SPLG). The second knock-out unit (HCKO2) will separateout CO2 from the recovered light gases via a 1-stage Rectisol unit(CO2SEP) for sequestration or recycle back to additional processunits (MXCO2C). The CO2 lean light gases will be recycled back tothe process.

The crude liquid hydrocarbons recovered from the two knock-out units is sent to a deethanizer (DEETH) to remove any C1-C2
hydrocarbons. The light HC gases are sent to an absorber column(ABS-COL) where a lean oil recycle is used to strip the C3+ HCsfrom the input. The liquid bottoms from the ABS-COL are thenrefluxed back to the deethanizer. The C3+ HCs from the bottom of

Chem

ttp(dcfaLbttwgaLfsfltduRmwis

2

bfCLu

T

wtbeAimtoDarpkeiDiLc1

sIaca

R.C. Baliban et al. / Computers and

he deethanizer are sent to a stabilizer column (STA-COL) wherehe C3/C4 hydrocarbons are removed and alkylated (ALK-UN) toroduce iso-octane and an LPG byproduct. Additional iso-butaneINBUT) may be fed to the alkylation unit for increased alkylate pro-uction. The bottoms from the stabilizer column is sent to a splitterolumn (SP-COL) to recover a lean oil recycle from the column topor use in the absorber column. Light and heavy gasoline fractionsre recovered from the column top and bottom, respectively. ThePG/alkylate from the alkylation unit is split (LPG-ALK) into an LPGyproduct (OUTLPG) and an alkylate fraction which is blended withhe gasoline fractions from the splitter column (OUTGAS). Each ofhe distillation units is modeled mathematically as a splitter unithere the split fraction of each species to an output stream is

iven by the information in the Process Flow Diagrams P850-A1501nd P850-A1502 from the NREL study (National Renewable Energyaboratory, 2011). All low pressure steam and cooling water neededor each of the units is derived for each of the units in the NRELtudy. The total amount of process utility that is needed per unitow rate from the top or bottom of the column is calculated, andhis ratio is used as a parameter in the process synthesis model toetermine the actual amount of each utility needed based on thenit flow rate. The alkylate was modeled as iso-butane (Nationalenewable Energy Laboratory, 2011) and the alkylation unit wasodeled using a species balance where the key species, butene,as completely converted to iso-butane. Butene is used as the lim-

ting species in this reaction because it is generally present in a farmaller concentration than iso-butane.

.6. Unit costs

The total direct costs, TDC, for the CBGTL refinery hydrocar-on production and upgrading units are calculated using estimatesrom several literature sources (Mobil Research & Developmentorporation, 1978, 1983, 1985; National Energy Technologyaboratory, 2007; National Renewable Energy Laboratory, 2011)sing the cost parameters in Table 3 and Eq. (4)

DC = (1 + BOP) · Co · S

So

sf

(4)

here Co is the installed unit cost, So is the base capacity, Sr ishe actual capacity, sf is the cost scaling factor, and BOP is thealance of plant (BOP) percentage (site preparation, utility plants,tc.). The BOP is estimated to be 20% of the total installed unit cost.ll numbers are converted to 2009 dollars using the GDP inflation

ndex (US Government Printing Office, 2009). Detailed cost esti-ates were not available for the MTO or OGD process units, so

he cost associated with these units was estimated from the costf an atmospheric MTG unit provided by Mobil (Mobil Research &evelopment Corporation, 1978). Note that not all units in Figs. 1–5re represented in Table 3. Some of the units shown in Table 3epresent the cost of that unit plus any auxillary units needed forroper unit operation. Specifically, (a) the three FT aqueous phasenock-out units are included in the cost of the hydrocarbon recov-ry column (Bechtel, 1998), (b) the cost of the FT ZSM-5 fractionators included in the cost of the FT ZSM-5 unit (Mobil Research &evelopment Corporation, 1983, 1985), (c) the MTO fractionator

s included in the cost of the MTO unit (National Renewable Energyaboratory, 2011), and (d) the OGD fractionator was included in theost of the OGD unit (Mobil Research & Development Corporation,978).

The total overnight capital, TOC, for each unit is calculated as theum of the total direct capital, TDC, plus the indirect costs, IC. The
C include engineering, startup, spares, royalties, and contingenciesnd is estimated to 32% of the TDC. The TOC for each unit must beonverted to a levelized cost to compare with the variable feedstocknd operational costs for the process. Using the methodology of
ical Engineering 47 (2012) 29– 56 37

Kreutz et al. (2008), the capital charges (CC) for the refinery are cal-culated by multiplying the levelized capital charge rate (LCCR) andthe interest during construction factor (IDCF) by the total overnightcapital (Eq. (5)).

CC = LCCR · IDCF · TOC (5)

Kreutz et al. (2008) calculates an LCCR value of 14.38%/year andIDCF of 1.076. Thus, a multiplier of 15.41%/year is used to convertthe overnight capital into a capital charge rate. Assuming an oper-ating capacity (CAP) of 330 days/year and operation/maintenance(OM) costs equal to 5% of the TOC, the total levelized cost (CostU)associated with a unit is given in Eq. (6).

CostUu =

(LCCR · IDCF

CAP+ OM

365

)·(

TOCu

Prod · LHVProd

)(6)

The levelized costs for the units described for hydrocarbon pro-duction and upgrading are added to the complete list of CBGTLprocess units given in Baliban, Elia, and Floudas (2012).

2.7. Objective function

The objective function for the model is given in Eq. (7). Thesummation represents the total cost of liquid fuels production andincludes contributions from the feedstocks cost (CostF), the elec-tricity cost (CostEl), the CO2 sequestration cost (CostSeq), and thelevelized unit investment cost (CostU). Each of the terms in Eq. (7)is normalized to the total lower heating value in GJ of productsproduced. For each case study, the capacity and ratio of liquid fuelproducts is fixed, so the normalization denominator in Eq. (7) willbe a constant parameter. Note that other objective functions (e.g.,maximizing the net present value) can be easily incorporated intothe model framework.

MIN∑u∈UIn

∑(u,s)∈SU

CostFs + CostEl + CostSeq +

∑u∈UInv

CostUu (7)

The process synthesis model with simultaneous heat, power,and water integration represents a large-scale non-convex mixed-integer non-linear optimization (MINLP) model that was solved toglobal optimality using a branch-and-bound global optimizationframework that was previously described (Baliban, Elia, Misener,et al., 2012). The MINLP model contains 32 binary variables, 11,104continuous variables, 10,103 constraints, and 351 non-convexterms (i.e., 285 bilinear terms, 1 trilinear term, 1 quadrilinear term,and 64 power functions). At each node in the branch-and-boundtree, a mixed-integer linear relaxation of the mathematical modelis solved using CPLEX (CPLEX, 2009) and then the node is branchedto create two children nodes. The solution pool feature of CPLEXis utilized during the solution of the relaxed model to generate aset of distinct points (150 for the root node and 10 for all othernodes), each of which is used as a candidate starting point to solvethe original model. For each starting point, the current binary vari-able values are fixed and the resulting NLP is minimized usingCONOPT. If the solution to the NLP is less than the current upperbound, then the upper bound is replaced with the NLP solutionvalue. At each step, all nodes that have a lower bound that is withinan � tolerance of the current upper bound ((LBnode/UB) ≥ 1 − �) areeliminated from the tree. For a more complete coverage of branch-and-bound algorithms, the reader is directed to the textbooks ofFloudas (Floudas, 1995, 2000) and reviews of global optimizationmethods (Floudas, Akrotirianakis, Caratzoulas, Meyer, & Kallrath,2005; Floudas & Gounaris, 2009; Floudas & Pardalos, 1995).

3. Computational studies

The proposed process synthesis model was used to ana-lyze twenty-four distinct case studies using perennial biomass


Table 3CBGTL refinery upgrading unit reference capacities, costs (2009$), and scaling factors.

Description Co (MM$) So SMax Units Scale basis sf Ref.

Fischer–Tropsch unit $12.26 23.79 60.0 kg/s Feed 0.72 b,c

Hydrocarbon recovery column $0.65 1.82 25.20 kg/s Feed 0.70 d

Distillate hydrotreater $2.25 0.36 81.90 kg/s Feed 0.60 d

Kerosene hydrotreater $2.25 0.36 81.90 kg/s Feed 0.60 d

Naphtha hydrotreater $0.68 0.26 81.90 kg/s Feed 0.65 d

Wax hydrocracker $8.42 1.13 72.45 kg/s Feed 0.55 d

Naphtha reformer $4.70 0.43 94.50 kg/s Feed 0.60 d

C5–C6 isomerizer $0.86 0.15 31.50 kg/s Feed 0.62 d

C4 isomerizer $9.50 6.21 – kg/s Feed 0.60 d

C3–C5 alkylation unit $52.29 12.64 – kg/s Feed 0.60 d

Saturated gas plant $7.83 4.23 – kg/s Feed 0.60 d

FT ZSM-5 reactor $4.93 10.60 – kg/s Feed 0.65 b,c

Methanol synthesis $8.22 35.647 – kg/s Feed 0.65 e

Methanol degasser $3.82 11.169 – kg/s Feed 0.70 e

Methanol-to-gasoline unit $5.80 10.60 – kg/s Feed 0.65 a,e

Methanol-to-olefins unit $3.48 10.60 – kg/s Feed 0.65 a

Olefins-to-gasoline/diesel unit $3.48 10.60 – kg/s Feed 0.65 a

CO2 separation unit $5.39 8.54 – kg/s Feed 0.62 a

Deethanizer $0.58 5.13 – kg/s Feed 0.68 a,e

Absorber column $0.91 0.96 – kg/s Feed 0.68 a,e

Stabilizer column $1.03 4.57 – kg/s Feed 0.68 a,e

Splitter column $1.01 3.96 – kg/s Feed 0.68 a,e

HF alkylation unit $8.99 0.61 – kg/s Feed 0.65 a,e

LPG/alkylate splitter $1.06 0.61 – kg/s Feed 0.68 a,e

a Mobil Research and Development Corporation (1978).b Mobil Research and Development Corporation (1983).

(ufthmtaeccctwowtbota(setecpcGr

lswso

c Mobil Research and Development Corporation (1985).d Bechtel Corporation (1998).e National Renewable Energy Laboratory (2011).

switchgrass), low-volatile bituminous coal (Illinois #6), and nat-ral gas as feedstocks. A global optimization framework was usedor each case study, and termination was reached if all nodes inhe branch-and-bound tree have been processed or if 100 CPUours have passed (Baliban, Elia, Misener, et al., 2012). The ulti-ate and proximate analysis of the biomass and coal feedstocks and

he molar composition of the natural gas feedstock are presenteds Supplementary Information. To examine the effects of potentialconomies of scale on the final liquid fuels price, three distinct plantapacities were examined to represent a small, medium, or largeapacity hybrid energy plant. Based on current petroleum refineryapacities (Energy Information Administration, 2009), representa-ive sizes of 10 thousand barrels per day (TBD), 50 TBD, and 200 TBDere chosen, respectively. A minimal carbon conversion threshold

f 40% was enforced for all of the case studies, and no upper boundas used for the amount of CO2 that is vented or sequestered. This

hreshold value was imposed to provide a comparative baselineetween all of the case studies, and does not have an effect on theverall process topologies. If no lower threshold value is imposed,hen the overall conversion for each study will range between 34%nd 39%, which is consistent with the results of a previous studyBaliban, Elia, Misener, et al., 2012). In general, raising the conver-ion rate produce more liquid fuels and decrease the byproductlectricity output from the plant, and for a more in-depth analysis,he reader is directed to the previous study (Baliban, Elia, Misener,t al., 2012). The overall greenhouse gas emission target for eachase study is set to have a 50% reduction from petroleum basedrocesses (Baliban, Elia, & Floudas, 2012; Baliban et al., 2011). Theurrent case studies do not include the cost of a carbon tax for anyHG emissions, though the process synthesis framework could be

eadily extended include a cost for the total lifecycle emissions.Four superstructure combinations will be investigated to ana-

yze the effect of plant topology on the final liquid fuels cost. These
uperstructures will consider (1) only Fischer–Tropsch synthesisith fractionation of the vapor effluent, (2) only Fischer–Tropsch
ynthesis with ZSM-5 catalytic upgrading of the vapor effluent, (3)nly methanol synthesis with either the MTG or MOGD process,

and (4) a comprehensive superstructure allowing all possibilitiesfrom (1), (2), or (3). Note that in superstructures (1), (2), and (4),any wax effluent from the Fischer–Tropsch units will be convertedto naphtha and diesel via a wax hydrocracker. Two sets of liquidfuels products (i.e., gasoline/diesel/kerosene and gasoline/diesel)will be considered to determine the effect of these products on theoptimal plant topology and overall costs. The ratio of liquid fuelproduction will be equal to the total 2010 United States demand(Energy Information Administration, 2011). Note that the processsuperstructure is also capable of analyzing a variable concentra-tion of output fuels (e.g., max diesel). Each of the 24 case studiesdiscussed below has a label P–CN where P is the type of productsproduced (GDK – gasoline/diesel/kerosene, GD – gasoline/diesel),C is the plant capacity (S – small, M – medium, L – large), and N isthe superstructure number defined above.

The cost parameters (Baliban, Elia, & Floudas, 2012; Balibanet al., 2011) used for CBGTL process are listed in Table 4. The costs forfeedstocks (i.e., coal, biomass, natural gas, freshwater, and butanes)include all costs associated with delivery to the plant gate. Theproducts (i.e., electricity and propane) are assumed to be sold fromthe plant gate, and do not include the costs expected for transportto the end consumer. The cost of CO2 capture and compression willbe included in the investment cost of the CBGTL refinery while thecost for transportation, storage, and monitoring of the CO2 is shownin Table 4.

Once the global optimization algorithm has completed, theresulting process topology provides (i) the operating conditionsand working fluid flow rates of the heat engines, (ii) the amountof electricity produced by the engines, (iii) the amount of cool-ing water needed for the engines, and (v) the location of the pinchpoints denoting the distinct subnetworks. Given this information,the minimum number of heat exchanger matches necessary tomeet specifications (i), (ii), (iii), and (iv) are calculated as pre-
viously described (Baliban, Elia, & Floudas, 2012; Baliban et al.,2011; Floudas, 1995; Floudas, Ciric, & Grossmann, 1986). Uponsolution of the minimum matches model, the heat exchanger topol-ogy with the minimum annualized cost can be found using the


Table 4Cost parameters (2009$) for the CBGTL refinery.

Item Cost Item Cost

Coal (LV bituminous) $93.41/short ton ($3/GJ) Biomass (Switchgrass) $139.97/dry metric ton ($8/GJ)Natural gas $5.39/TSCF1 ($5.5/GJ) Freshwater $0.50/metric tonButanes $1.84/gallon Propanes $1.78/gallon

2

sFas

3

fp1twiivctpibtlFss

tirghvaeltivbtirt

9isptsvuiHd

Electricity $0.07/kWh

1 TSCF – thousand standard cubic feet2 TS&M – transportation, storage, and monitoring

uperstructure methodology (Elia et al., 2010; Floudas, 1995;loudas et al., 1986). The investment cost of the heat exchangers isdded to the investment cost calculated within the process synthe-is model to obtain the final investment cost for the superstructure.

.1. Optimal process topologies

The information detailing the optimal process topologiesor all case studies is shown in Table 5. Three possible tem-erature options were used for the biomass gasifier (900 ◦C,000 ◦C, 1100 ◦C), the coal gasifier (1100 ◦C, 1200 ◦C, 1300 ◦C),he auto-thermal reactor (700 ◦C, 800 ◦C, 950 ◦C), and the reverseater–gas-shift unit (400 ◦C, 500 ◦C, 600 ◦C). For all 24 case stud-

es, the biomass and coal solid/vapor fueled gasifiers were utilizedn the optimal process design. Thus, each gasifier employed aapor phase recycle stream as a fuel input along with the solidoal or biomass. Recycle of some of the unreacted synthesis gaso the gasifiers helped to consume some CO2 generated in therocess and reduce the overall process emissions by convert-

ng the CO2 to CO for additional liquid fuels production. For theiomass gasifier, the 900 ◦C unit is always selected for superstruc-ure 1 and only selected for superstructure 3 at high capacityevels. For all other case studies, the 1100 ◦C unit is selected.or the coal gasifier, the 1300 ◦C unit was always selected foruperstructures 1, 2, and 3 and the 1100 ◦C unit was selected foruperstructure 4.

Selection of the gasifier operating temperatures in the optimalopology represents a balance between (i) the levels of oxidantnput to the gasifier, (ii) the extent of consumption of CO2 via theeverse water–gas-shift reaction, and (iii) the level of waste heatenerated from syngas cooling. Lower gasifier temperatures willave less favorable conditions for CO2 consumption due to loweralues of the water–gas-shift equilibrium constant and a smallermount of waste heat for use in steam generation and ultimatelylectricity production. However, lower temperatures will requireower levels of O2 for combustion within the gasifier which reduceshe investment and utility cost for oxygen generation and mayncrease the overall efficiency of the gasifier. The alternative disad-antages with a higher O2 in the higher temperature gasifiers arealanced by an increase in the CO2 reduction potential and the addi-ional waste-heat generated. The operating temperature selectedn the 24 case studies reflects the trade-offs between emissionseduction, electricity production, and overall process efficiency forhe entire refinery.

The auto-thermal reformer temperature was selected to be50 ◦C for twelve of the case studies and 800 ◦C for the remain-

ng twelve studies (see Table 5). A 950 ◦C unit is always used foruperstructure 1, used for superstructure 2 in the medium and largelants, and used in superstructure 4 for the large plants. Selec-ion of the temperature for the auto-thermal reformer will haveimilar topological effects as the gasifiers, though the overall con-ersion of CH4 will also increase with increasing temperature. The
se of the highest temperature reformer is beneficial since approx-
mately 90% of the input CH4 can be converted to syngas using a2O/CH4 ratio of approximately 1.2–1.5. Ultimately, this will alsoecrease the working capacity of the FT synthesis or methanol

CO2 TS&M $10/metric ton

synthesis units because the input CH4 is an inert species that willnot be separated until downstream of these units. The selectionof the 800◦C units for the remaining studies generally converts82–85% of the CH4, though the decrease in the oxygen require-ment to the unit provides an economic benefit to the decreasedconversion of the natural gas.

A dedicated reverse water–gas-shift unit was not selected foreither product composition and plant capacity that used super-structures 1, 2, or 3. For each of these case studies, the proper syngasratio requirements for the FT and methanol synthesis was met vialight gas recycle to either the gasifiers or the auto-thermal reactorunits. For the case studies using superstructure 4, a 600 ◦C reversewater–gas-shift unit was utilized to both consume CO2 generatedin the process and shift the syngas ratios for conversion. All of thecase studies generated H2 using pressure-swing absorption and O2using air separation. The H2 was utilized mostly for product upgrad-ing and for injection, with the balance being sent to the reversewater–gas-shift units to consume some CO2. Note that H2 sepa-ration is required for hydrotreating and hydrocracking within theproduct upgrading section. Electrolyzers were not utilized in anycase study due to the high capital ($500/kW) and electricity costsof the unit. The electricity input to the electrolyzers is assumedto come from a non-carbon based source (e.g., wind/solar), whichwas assumed to have a high cost (i.e., $0.10/kWh). Note that inputelectricity from a carbon-based source (i.e., biomass/coal/naturalgas) is not considered because the process superstructure accountsfor H2 generation from pressure-swing absorption. A decrease inthe non-carbon based electricity cost may have an effect on theelectrolyzer use, as noted in a previous study (Baliban et al., 2011).Both a gas and steam turbine are used in each case study to pro-duce electricity for the process and to partially sell as a byproduct.To reduce the GHG emissions from the processes, each case studyutilized CO2 capture and sequestration both upstream of synthesisgas conversion and downstream of the gas turbine engine.

The case studies using superstructures 1 and 2 required FTsynthesis of the hydrocarbons, and each case study utilized aniron-based catalyst within both the minimal-wax and nominal-wax reactors. Additionally, the reverse water–gas-shift reactionwas facilitated in most of the case studies, with the exception ofthe minimal-wax reactor in superstructure 2 for the medium andlarge capacities. In the former case studies, the iron-based unitscan take advantage of the exothermic FT reaction to provide heatfor the endothermic reverse water–gas-shift reaction (Baliban, Elia,& Floudas, 2012; Baliban et al., 2011). In the latter studies, theadditional CO2 that is generated from the FT reactors is capturedand recycled back to the process to minimize the GHG emissions.Due to the constraints of the process superstructure, upgradingof the vapor phase FT effluent utilized a fractionation scheme forsuperstructure 1 and the ZSM-5 catalyst for superstructure 2. Forsuperstructure 3, the syngas was converted to methanol ratherthan hydrocarbons via the FT reaction. For all case studies usingthis superstructure, both the methanol-to-gasoline and methanol-
to-olefins/distillate processes are utilized to produce the liquidfuels in the appropriate output ratios. In the case studies usingsuperstructure 4, the technologies used for liquid fuels productionare highly dependent on the plant capacity and the type of fuels


Table 5Topological information for the optimal solutions for the 24 case studies. Specifically listed is the operating temperature of the biomass gasifier (BGS), the coal gasifier (CGS),the auto-thermal reactor (ATR), and the reverse water–gas-shift unit (RGS). The gasifiers are also labeled as either solid/vapor (S/V) or solid (S) fueled, implying the presenceor absence of vapor-phase recycle process streams. The presence of a CO2 sequestration system (CO2SEQ) or a gas turbine (GT) is noted using yes (Y) or no (-). The minimumwax 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 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 (-). The results for the completesuperstructure and medium sized capacity (M4) are shown in boldface.

Case study GDK-S1 GDK-S2 GDK-S3 GDK-S4 GDK-M1 GDK-M2 GDK-M3 GDK-M4 GDK-L1 GDK-L2 GDK-L3 GDK-L4

BGS Temp. (◦C) 900 1100 1100 1100 900 1100 1100 1100 900 1100 900 1100BGS Type S/V S/V S/V S/V S/V S/V S/V S/V S/V S/V S/V S/VCGS Temp. (◦C) 1300 1300 1300 1100 1300 1300 1300 1100 1300 1300 1300 1100CGS Type S/V S/V S/V S/V S/V S/V S/V S/V S/V S/V S/V S/VRGS Temp. (◦C) – – – 600 – – – 600 – – – 600ATR Temp. (◦C) 950 800 800 800 950 950 800 800 950 950 800 950Min Wax FT Ir. rWGS Ir. rWGS – – Ir. rWGS Ir. fWGS – – Ir. rWGS Ir. fWGS – –Nom. Wax FT Ir. rWGS Ir. rWGS – Ir. rWGS Ir. rWGS Ir. rWGS – Ir. rWGS Ir. rWGS Ir. rWGS – –FT Upgrading Fract. ZSM-5 – Fract. Fract. ZSM-5 – Fract. Fract. ZSM-5 – –MTG Usage – – Y Y – – Y Y – – Y YMOGD Usage – – Y Y – – Y Y – – Y YCO2SEQ Usage Y Y Y Y Y Y Y Y Y Y Y YGT Usage Y Y Y Y Y Y Y Y Y Y Y Y

Case study GD-S1 GD-S2 GD-S3 GD-S4 GD-M1 GD-M2 GD-M3 GD-M4 GD-L1 GD-L2 GD-L3 GD-L4

BGS Temp. (◦C) 900 1100 1100 1100 900 1100 1100 1100 900 1100 900 1100BGS type S/V S/V S/V S/V S/V S/V S/V S/V S/V S/V S/V S/VCGS Temp. (◦C) 1300 1300 1300 1100 1300 1300 1300 1100 1300 1300 1300 1100CGS type S/V S/V S/V S/V S/V S/V S/V S/V S/V S/V S/V S/VRGS Temp. (◦C) – – – 600 – – – 600 – – – 600ATR Temp. (◦C) 950 800 800 800 950 950 800 800 950 950 800 950Min wax FT Ir. rWGS Ir. rWGS – – Ir. rWGS Ir. rWGS – – Ir. rWGS Ir. rWGS – –Nom. wax FT Ir. rWGS Ir. rWGS – Ir. rWGS Ir. rWGS Ir. rWGS – Ir. rWGS Ir. rWGS Ir. rWGS – Ir. rWGSFT upgrading Fract. ZSM-5 – Fract. Fract. ZSM-5 – Fract. Fract. ZSM-5 – ZSM-5MTG usage – – Y Y – – Y Y – – Y YMOGD usage – – Y – – – Y – – – Y –

Y

Y

pwpFactwGn

ccuoaahb

3

c(bo(m(2

CO2SEQ usage Y Y Y Y Y

GT usage Y Y Y Y Y

roduced. For the six studies with superstructure 4, the minimal-ax FT unit was never utilized and the methanol-to-gasolinerocess was always utilized. The nominal-wax iron-based rWGST unit was used for the two small plants, the two medium plants,nd the large plant that does not produce kerosene. For the fivease studies that used FT, the vapor phase was always convertedo gasoline-range hydrocarbons using ZSM-5. The MOGD processas used to generate diesel and kerosene for all plant sizes in theDK case studies. In the GD case studies, the MOGD process wasot utilized and all diesel was generated from wax hydrocracking.

The results for the complete superstructure and medium sizedapacity (M4) are shown in boldface in Table 5. For each of theseases, both the biomass and coal gasifiers were solid/vapor fuelednits operating at 1100 ◦C. A dedicated reverse water–gas-shift unitperating at 600 ◦C is used and the auto-thermal reactor operatest 800 ◦C for both studies. The liquid fuels are produced via (i) cat-lytic ZSM-5 upgrading of the iron-based rWGS FT effluent, (ii) waxydrocracking, and (iii) methanol-to-gasoline for both studies andy MOGD for the study requiring kerosene production.

.2. Overall costs of liquid fuels

The overall cost of liquid fuel production (in $/GJ) is based on theosts of feedstocks, capital investment, operation and maintenanceO&M), and CO2 sequestration and can be partially defrayed usingyproduct sales of LPG and electricity. Feedstock costs are basedn the as-delivered price for (i) the three major carbon feedstocks
coal, biomass, and natural gas), (ii) butanes needed for the iso-
erization process (Baliban et al., 2010, 2011; Bechtel, 1992), andiii) freshwater needed to make-up for process losses(Baliban et al.,012a). Table 6 outlines the breakdown of the cost contribution for

Y Y Y Y Y YY Y Y Y Y Y

each case study, as well as the lower bound and the optimality gapvalues. The total cost is also converted into a break-even oil price(BEOP) in $/barrel based on the refiner’s margin for gasoline, diesel,or kerosene (Baliban et al., 2011; Kreutz et al., 2008), and repre-sents the price of crude oil at which the CBGTL process becomeseconomically competitive with petroleum based processes.

The overall cost values range between $17.33 and $18.79/GJfor a small plant, $16.06–$17.66/GJ for a medium plant, and$14.76–$16.20/GJ for a large plant. For a medium sized plant pro-ducing gasoline, diesel, and kerosene, the optimization model forthe complete superstructure (i.e., case study GDK-M4) selects atopology with an overall cost of $16.25/GJ or $79.83/bbl crude oilequivalent. The upper bound value found at the termination ofthe global optimization algorithm is 4.56% above the lower boundvalue of $15.51/GJ. When only gasoline and diesel are producedin the general medium sized plant (GD-M4), the overall cost ofliquid fuel production for a medium sized plant with the mostgeneral superstructure is $16.06/GJ or $78.74/bbl crude oil equiv-alent with a 5.35% optimality gap from its lower bound value of$15.20/GJ. Negative values in the cost contributions from elec-tricity and propane represent the profit gained from selling theseitems as byproducts. In all of the 24 case studies, the selectedplant topologies are net producers of electricity and propane(see Table 6).

For a given capacity level, Table 6 shows that the lowest overallcost is achieved through the use of the most general superstruc-ture topology. Additionally, the second lowest cost is consistently
found with superstructure 3, suggesting that the methanol synthe-sis/conversion process units generally yield a plant design with alower overall cost. However, the decrease in cost between super-structure 3 (only methanol) and superstructure 4 (methanol/FT)


Table 6Overall cost results for the 24 case studies. The case studies where the plant topologies produce gasoline, diesel, and kerosene are labeled as GDK, and the topologies thatproduce gasoline and diesel are labeled as GD. The small (S), medium (M), and large (L) case studies are each labeled with the superstructure number, where (1) indicates thatonly Fischer–Tropsch synthesis with fractionation of the vapor effluent is considered, (2) only Fischer–Tropsch synthesis with ZSM-5 catalytic upgrading of the vapor effluent,(3) only methanol synthesis with either the MTG or MOGD process, and (4) a comprehensive superstructure allowing all possibilities from (1), (2), or (3). The contributionto the total costs (in $/GJ) come from coal, biomass, natural gas, butanes, water, CO2 sequestration (CO2. Seq.), and the investment. Propane is always sold as a byproductwhile electricity may be sold as a byproduct (negative value). The overall costs are reported in ($/GJ) and ($/bbl) basis, along with the lower bound values in ($/GJ) and theoptimality gap between the reported solution and the lower bound. The results for the complete superstructure and medium sized capacity (M4) are shown in boldface.

Contribution to cost($/GJ of products)

Case study

GDK-S1 GDK-S2 GDK-S3 GDK-S4 GDK-M1 GDK-M2 GDK-M3 GDK-M4 GDK-L1 GDK-L2 GDK-L3 GDK-L4

Coal 3.15 3.18 3.20 2.91 3.32 3.31 3.05 3.16 3.21 3.08 3.33 3.37Biomass 2.71 2.69 2.70 2.74 2.75 2.81 2.73 2.70 2.69 2.73 2.69 2.67Natural gas 3.58 3.48 3.43 4.14 3.08 3.02 3.80 3.53 3.40 3.78 3.14 3.02Butane 0.28 0.31 0.40 0.28 0.29 0.25 0.34 0.33 0.36 0.25 0.34 0.36Water 0.03 0.02 0.03 0.02 0.02 0.02 0.02 0.02 0.02 0.02 0.02 0.02CO2 Seq. 0.51 0.50 0.51 0.50 0.50 0.49 0.51 0.50 0.51 0.51 0.51 0.51Investment 11.15 10.81 10.22 10.03 8.29 8.16 7.50 7.65 7.25 7.44 6.64 6.70O&M 3.27 3.17 3.00 2.94 2.43 2.40 2.20 2.25 2.13 2.18 1.95 1.97Electricity −5.69 −5.43 −5.26 −5.96 −2.86 −2.82 −3.34 −3.72 −3.20 −3.92 −3.02 −3.48Propane −0.19 −0.15 −0.20 −0.15 −0.17 −0.14 −0.20 −0.17 −0.17 −0.21 −0.22 −0.19Total ($/GJ) 18.79 18.59 18.02 17.46 17.66 17.51 16.61 16.25 16.20 15.86 15.37 14.95Total ($/bbl) 94.32 93.18 89.90 86.72 87.85 87.00 81.85 79.83 79.52 77.58 74.84 72.40Lower bound ($/GJ) 17.73 17.86 17.31 16.92 16.68 16.54 15.83 15.51 15.35 15.32 14.74 14.40Gap 5.63% 3.92% 3.92% 3.10% 5.52% 5.55% 4.67% 4.56% 5.24% 3.40% 4.16% 3.67%

Contribution to cost($/GJ of products)

Case study

GD-S1 GD-S2 GD-S3 GD-S4 GD-M1 GD-M2 GD-M3 GD-M4 GD-L1 GD-L2 GD-L3 GD-L4

Coal 2.71 3.38 2.75 2.72 3.25 3.13 3.34 3.23 3.19 3.39 3.27 3.27Biomass 2.75 2.65 2.74 2.75 2.68 2.68 2.66 2.66 2.68 2.65 2.66 2.65Natural gas 4.62 2.98 4.51 4.59 3.30 3.56 3.07 3.30 3.42 2.95 3.24 3.21Butane 0.26 0.26 0.31 0.32 0.30 0.38 0.33 0.36 0.33 0.27 0.29 0.33Water 0.03 0.03 0.02 0.02 0.02 0.03 0.02 0.02 0.02 0.02 0.03 0.03CO2 Seq. 0.50 0.50 0.50 0.50 0.50 0.50 0.50 0.50 0.50 0.50 0.50 0.50Investment 11.28 11.09 10.22 10.07 8.11 8.38 7.33 7.48 6.85 6.99 6.33 6.52O&M 3.31 3.26 3.00 2.96 2.38 2.46 2.15 2.20 2.01 2.05 1.86 1.91Electricity −6.53 −5.72 −5.91 −6.43 −3.01 −3.65 −2.84 −3.56 −2.83 −2.81 −2.73 −3.53Propane −0.16 −0.20 −0.16 −0.16 −0.21 −0.17 −0.20 −0.14 −0.18 −0.20 −0.14 −0.14Total ($/GJ) 18.59 18.06 17.99 17.33 17.33 17.30 16.36 16.06 16.01 15.84 15.30 14.76Total ($/bbl) 93.16 90.14 89.73 85.99 85.99 85.82 80.46 78.74 78.45 77.47 74.42 71.31

%

itiosusGZsttfa

3

fccrcnu

Lower bound ($/GJ) 17.77 17.28 17.17 16.41 16.52Gap 4.39% 4.30% 4.55% 5.32% 4.67

mplies that there is a degree of synergy that can be achievedhrough the use of both technologies. The resulting level of synergys likely to be tied to the capacity of the plant and the compositionf liquid fuels that will be produced. The CBGTL case studies usinguperstructures 2 (FT with ZSM-5 upgrading) have a lower costltimately due to a decrease in the complexity of the FT synthe-is and upgrading section of the plant. In some case studies (i.e.,DK-L2, GD-L2, and GD-M2), the investment cost of the plant withSM-5 upgrading was higher than that for the corresponding casetudy without ZSM-5 upgrading. The increase in investment is dueo a higher overall flow rate of syngas through the refinery dueo (i) increased recycle flow of the unreacted syngas to decreaseeedstock costs or (2) increased flow of the feedstocks to producedditional byproduct electricity.

.3. Parametric analysis

Table 6 indicates that the largest contribution to the overalluels cost is associated with the capital investment (i.e., capitalharges and operation/maintenance). A reduction in total plantost may be achieved through innovation of novel technologies
ather than relying on economies of scale for more mature pro-esses (Adams & Barton, 2011). However, the coal, biomass, andatural gas may have a wide variability in the overall cost of liq-id fuel production. Depending on the demand for these materials
16.40 15.86 15.20 15.10 14.99 14.79 13.975.20% 3.07% 5.35% 5.67% 5.34% 3.33% 5.33%

and the plant location throughout the country, the feedstock costsmay be higher or lower than the national average. Given the deliv-ered feedstock costs in Table 4 and the feedstock lower heatingvalues in Supplementary Table S1, the cost per unit energy is calcu-lated for coal ($3.0/GJ), biomass ($8.0/GJ), and natural gas ($5.5/GJ).These cost parameters represent conservative estimates (EnergyInformation Administration, 2011; Kreutz et al., 2008; Larson et al.,2009; National Academy of Sciences, 2009) for the total deliveredcost of a particular feedstock, and it is important to investigate howthe BEOP will be affected if these cost parameters are reduced. Asan illustrative example, the BEOP for case study GDK-M4 is calcu-lated assuming either low, nominal, or high cost values for each ofthe three feedstocks. These respective values are (i) $2/GJ, $2.5/GJ,and $3/GJ for coal, (ii) $5/GJ, $6.5/GJ, and $8/GJ for biomass, and (iii)$4/GJ, $4.75/GJ, and $5.5/GJ for natural gas. The BEOP was calculatedfor each of the 27 parameter combinations, and the histogram ofresults is shown in Fig. 6.

Each cost bin in Fig. 6 represents a $2/barrel window for theBEOP. That is, the first bin represents all of the parameter combina-tions that had a BEOP between $60/bbl and $62/bbl, the secondbin is between $62/bbl and $64/bbl, and so on. The histogram
shows a Gaussian-like distribution with two major peaks in the$68/bbl–$72/bbl range with a total of 13 counts. The shape of thehistogram can be inferred from Table 6 since the contribution ofeach feedstock to the overall cost is relatively similar. The singular

42 R.C. Baliban et al. / Computers and Chem

0

1

2

3

4

5

6

7

8

60 64 68 72 76

Co

un

ts

BEOP Bin ($/bbl)

Fcv

poep

3

tthpgtgcgas

M$ft$pcGslrf

esrlscssGpt

ig. 6. Parametric analysis of feedstock cost. The histogram shows the number ofounts (out of 27) for break-even oil price (BEOP) when low, nominal, and highalues are used for the costs of coal, biomass, and natural gas.

eak in the leftmost bin corresponds to a BEOP of $62.7/bbl and isbtained if the low parameters are used for each feed. The high-st BEOP is equal to $80.0/bbl, and is obtained if all of the higharameter values are used.

.4. Investment costs

The plant investment cost is further decomposed into cost con-ributions from different sections of the plant in Table 7, namelyhe syngas generation, syngas cleaning, hydrocarbon production,ydrocarbon upgrading, hydrogen/oxygen production, heat andower integration, and wastewater treatment sections. The syngaseneration section is consistently the highest contributing factor inhe investment cost due to the capital intensive coal and biomassasifier units. The next highest contributing factors are the syngasleaning, hydrogen/oxygen production, and heat and power inte-ration sections, followed by the hydrocarbon production section,nd finally the hydrocarbon upgrading and wastewater treatmentections.

The total investment cost ranges from $1166 to $1296M for small plants producing gasoline, diesel, and kerosene,

4359–$4823 MM for medium plants, and $15,446–$17,309 MMor large plants. The normalized investment costs, however, revealhe economies of scale obtained in larger sized plants, ranging from116k to $130k/bpd for small plants, $87k–$96k/bpd for mediumlants, and $78k–$87k/bpd for large plants. Among the small plantase studies, the case with the most general superstructure (i.e.,DK-S4) is able to achieve the lowest investment cost. For largerized plants, however, GDK-M3 and GDK-L3 case studies have theowest investment costs for medium and large plants case studies,espectively. Conversely, the case studies using superstructure 1rom all capacity levels have the highest total investment cost.

Comparisons between the GDK and GD case studies reveal inter-sting trade-offs in investment costs. For the small plants casetudies, plant topologies that produce only gasoline and dieselesult in higher investment costs than the ones that produce gaso-ine, diesel, and kerosene. The increased cost of the small GD casetudies is due to a higher flow rate of syngas throughout the pro-ess units due to a slightly higher level of recycle than the GDKmall case studies. The increased investment costs for the small GD
tudies do lead into smaller levels of feedstock usage than the smallDK studies, and therefore have a lower overall cost of liquid fuelsroduction (see Table 6). For the medium and large GD case studies,he topologies that produce gasoline and diesel fuels consistently

yield lower total investment costs than their GDK counterpartsdue to the less complicated refining that is needed to producekerosene.

3.5. Material and energy balances

The overall material and energy balances for the 24 case stud-ies are shown in Tables 8 and 9, respectively. The biomass andcoal flow rates are based of dry tons (dt) while the natural gas isshown in million standard cubic feet (mscf). From Tables 8 and 9,it can be seen that coal provides the most energy input to theplant, followed generally by natural gas, and then biomass. Forexample, the most small capacity plant with the most generalsuperstructure (GDK-S4) requires 69.56 dt/h coal, 51.08 dt/h forbiomass, and 1.83 mscf/h natural gas. These values correspondto 596 MW energy input from coal, 224 MW from biomass, and497 MW from natural gas. This distribution remains relatively con-sistent when the plant size increases. For the medium sized plant(case study GDK-M4), 377.39 dt/h is needed for coal, 251.95 dt/hfor biomass, and 7.77 mscf/h, corresponding to 3234 MW energyinput for coal, 1106 MW for biomass, and 2114 MW for natural gas.Case study GDK-L4 requires 1607.23 dt/h coal, 997.60 dt/h biomass,and 26.64 mscf/h natural gas, corresponding to 13,775 MW energyinput from coal, 4377 MW from biomass, and 7250 MW from natu-ral gas. The smaller contribution of biomass relative to the othertwo feedstocks is due to the higher $/GJ costs associated withbiomass. The highest driving force for the use of biomass is thelifecycle GHG reduction potential, but the use of CO2 sequestra-tion from the 24 case studies (see Table 8) will reduce the biomassrequirement for the plant. A restriction on the amount of CO2 thatis captured for sequestration (e.g., no nearby available locationsfor CO2 storage) will ultimately increase the biomass feedstockrequirement, and the biomass could become the largest energycontributor to the refinery. The authors note that the biomassrequirement for the large case studies (i.e., 200,000/bpd) is nec-essary to achieve a life-cycle GHG emissions that is 50% lowerthan petroleum-based processes. Though the biomass-based plantdesigns by the National Renewable Energy Laboratory use approxi-mately 2000 dry tons/day (National Renewable Energy Laboratory,2011; Spath et al., 2005), the availability of biomass may be substan-tially higher in several counties (e.g., Midwestern United States)after land-use change or an increase in crop yields (Department ofEnergy, 2005).

Almost all of the case studies do not vent CO2 from the process,and utilize CO2 sequestration to reduce the lifecycle GHG emissionsof the plant. The GDK-M1 and GDK-M2 studies vent a small amountof CO2, though the CO2 is only 1–2% of the total CO2 produced bythe plant. The balance of the CO2 is captured for sequestration. Thehigh utilization of CO2 sequestration allows for an increased use ofthe cheaper fossil fuels coal and natural gas, which can be anywherefrom $3/GJ to $6/GJ less expensive than biomass. The biomass doesprovide negative emission values from CO2 intake from the atmo-sphere during cultivation and additional soil storage from land usechange, so a level of biomass input on a mass/energy basis that isroughly equivalent to that of coal or natural gas is still required.

The electricity production ranges from 179 to 221 MW for smallplants, 478–631 MW for medium plants, and 1850–2661 MW forlarge plants. In all case studies, a high amount of electricity is pro-duced to help lower the overall cost of fuels for the plant. Theelectricity output also improves the efficiency of the topologies,with GD-S2, GDK-S1, and GD-S1 achieving the highest energy effi-ciencies (i.e., 67.1%, 65.8%, and 65.7%, respectively) compared to
other case studies in their subcategories (see Table 9). The energyefficiency values are calculated by dividing the total energy out-put (i.e., fuel products, propane, or electricity) by the total energyinput (i.e., via coal, biomass, natural gas, butane, or electricity). If


Table 7Breakdown of the investment costs for the 24 case studies. The major sections of the plant include the syngas generation section, syngas cleaning, hydrocarbon production,hydrocarbon upgrading, hydrogen/oxygen production, heat and power integration, and wastewater treatment blocks. The values are reported in MM$ and normalized withthe amount of fuels produced ($/bpd). The results for the complete superstructure and medium sized capacity (M4) are shown in boldface.

Contribution to cost (MM$) Case study


Syngas generation 494 492 476 478 1422 1443 1314 1369 5362 5702 5063 5187Syngas cleaning 240 234 222 225 813 786 769 758 3153 3186 2932 2870Hydrocarbon production 218 208 170 166 738 731 566 603 2682 2775 1985 2148Hydrocarbon upgrading 24 22 16 16 165 147 96 100 343 326 207 206Hydrogen/oxygen production 145 138 139 126 789 770 768 754 2424 2495 2588 2451Heat and power integration 146 137 138 129 781 742 747 767 2521 2412 2364 2405Wastewater treatment 29 26 28 26 115 127 99 98 377 412 307 305Total (MM $) 1296 1258 1188 1166 4823 4745 4359 4450 16,862 17,309 15,446 15,572Total ($/bpd) 129,647 125,754 118,809 116,609 96,451 94,897 87,177 88,993 87,211 86,547 79,335 77,858

Contribution to cost (MM$) Case study


Syngas generation 500 511 486 483 1373 1480 1270 1318 5068 5183 4875 4868Syngas cleaning 244 242 219 223 785 805 779 764 3071 3035 2770 2919Hydrocarbon production 218 209 161 166 735 785 554 594 2399 2602 1867 2190Hydrocarbon upgrading 25 22 16 15 154 150 90 86 338 327 190 205Hydrogen/oxygen production 148 139 134 131 784 756 742 734 2365 2413 2471 2292Heat and power integration 146 140 143 126 767 773 733 759 2326 2304 2237 2381

147

94,3

etaeek

TOk(a

Wastewater treatment 30 27 30 27Total (MM $) 1311 1290 1189 1171

Total ($/bpd) 131,118 128,975 118,893 117,083

lectricity is output from the system, the value is listed as nega-ive in Table 8 and the magnitude of the energy value in Table 9 isdded to the total output. If the value is positive in Table 8, then this
nergy is added to the total input to the system. The overall energyfficiency of the CBGTL topologies producing gasoline, diesel, anderosene ranges between 58.5 and 67.1% for all plant sizes.
able 8verall material balance for the 24 case studies. The inputs to the CBGTL process are biomerosene, LPG, sequestered and vented CO2, and electricity. Biomass and coal are input inmscf/h), liquids in thousand barrels per day (kBD), and CO2 in metric tons per hour (tonnre shown in boldface.

Material balances Case study

GDK-S1 GDK-S2 GDK-S3 GDK-S4 GDK-M1 GDK

Biomass (dt/h) 50.56 50.28 50.42 51.08 256.34 26Coal (dt/h) 75.16 75.95 76.31 69.56 396.13 39Natural gas (mscf/h) 1.58 1.53 1.51 1.83 6.78

Butane (kBD) 0.21 0.23 0.30 0.21 1.07

Water (kBD) 18.18 14.34 18.85 16.01 80.05 7Gasoline (kBD) 6.72 6.72 6.72 6.72 33.60 3Diesel (kBD) 2.15 2.15 2.15 2.15 10.77 1Kerosene (kBD) 1.13 1.13 1.13 1.13 5.63

LPG (kBD) 0.14 0.11 0.15 0.11 0.63

Seq. CO2 (tonne/h) 240.04 239.65 240.14 239.61 1183.81 116Vented CO2 (tonne/h) 0.00 0.00 0.00 0.00 15.23 2Electricity (MW) −193.14 −184.16 −178.57 −202.29 −484.88 −47

Material balances Case study

GD-S1 GD-S2 GD-S3 GD-S4 GD-M1 GD

Biomass (dt/h) 51.33 49.47 51.21 51.32 249.75 2Coal (dt/h) 64.64 80.78 65.60 64.81 387.91 3Natural gas (mscf/h) 2.04 1.31 1.99 2.02 7.27

Butane (kBD) 0.19 0.20 0.23 0.24 1.11

Water (kBD) 20.51 19.18 12.51 14.84 79.22

Gasoline (kBD) 7.57 7.57 7.57 7.57 37.86

Diesel (kBD) 2.43 2.43 2.43 2.43 12.14

LPG (kBD) 0.12 0.15 0.12 0.12 0.79

Seq. CO2 (tonne/h) 238.72 239.15 238.76 238.80 1196.62 11Vented CO2 (tonne/h) 0.00 0.00 0.00 0.00 0.00

Electricity (MW) −221.38 −193.96 −200.46 −218.06 −509.61 −6

20 123 94 94 367 401 314 29817 4872 4261 4348 15,935 16,265 14,723 15,15335 97,434 85,226 86,958 79,677 81,326 73,615 75,764

3.6. Carbon and greenhouse gas balances

The overall carbon balance for the CBGTL processes is shown
in Table 10 and highlights the eight major points where carbon iseither input or output from the system. The results for the completesuperstructure and medium sized capacity (M4) case studies are
ass, coal, natural gas, butane, and water, while the outputs include gasoline, diesel, dry metric tons per hour (dt/h), natural gas in million standard cubic feet per houre/h). The results for the complete superstructure and medium sized capacity (M4)

-M2 GDK-M3 GDK-M4 GDK-L1 GDK-L2 GDK-L3 GDK-14

2.08 254.39 251.95 1005.02 1018.08 1002.86 997.605.36 364.37 377.39 1532.36 1468.83 1588.67 1607.236.66 8.38 7.77 29.98 33.28 27.65 26.640.92 1.26 1.21 5.28 3.71 5.00 5.387.55 75.88 68.38 296.85 333.54 306.86 313.533.60 33.60 33.60 134.39 134.39 134.39 134.390.77 10.77 10.77 43.10 43.10 43.10 43.105.63 5.63 5.63 22.51 22.51 22.51 22.510.51 0.74 0.62 2.53 3.14 3.25 2.857.69 1200.70 1198.94 4796.55 4805.53 4807.09 4801.239.69 0.00 0.00 0.00 0.00 0.00 0.007.54 −567.06 −631.42 −2171.37 −2661.25 −2045.52 −2357.99

-M2 GD-M3 GD-M4 GD-L1 GD-L2 GD-L3 GD-L4

50.28 248.04 248.07 998.77 988.91 991.27 990.1873.28 398.69 385.57 1523.11 1619.62 1558.90 1562.09

7.84 6.75 7.27 30.17 26.03 28.55 28.281.42 1.24 1.35 4.90 4.05 4.33 4.95

92.56 68.38 79.22 286.85 350.22 366.90 360.2337.86 37.86 37.86 151.44 151.44 151.44 151.4412.14 12.14 12.14 48.56 48.56 48.56 48.56

0.61 0.75 0.51 2.61 2.90 2.13 2.0894.09 1196.04 1192.62 4778.68 4782.98 4771.66 4770.88

0.00 0.00 0.00 0.00 0.00 0.00 0.0019.46 −481.53 −602.94 −1916.43 −1908.04 −1849.71 −2393.96


Table 9Overall energy balance for the 24 case studies. The energy inputs to the CBGTL process come from biomass, coal, natural gas, and butane, and the energy outputs are gasoline,diesel, kerosene, LPG, and electricity. The energy efficiency of the process is calculated by dividing the total energy output with the total energy inputs to the process.

Energy balances (MW) Case study


Biomass 222 221 221 224 1125 1150 1116 1106 4410 4467 4401 4377Coal 644 651 654 596 3395 3388 3123 3234 13,133 12,589 13,616 13,775Natural gas 429 417 411 497 1845 1812 2279 2114 8157 9057 7522 7250Butane 13 14 18 13 65 56 77 74 321 226 304 327Gasoline 428 428 428 428 2141 2141 2141 2141 8563 8563 8563 8563Diesel 153 153 153 153 766 766 766 766 3065 3065 3065 3065Kerosene 78 78 78 78 390 390 390 390 1558 1558 1558 1558LPG 9 7 9 7 38 31 45 38 154 191 197 173Electricity 193 184 179 202 485 478 567 631 2171 2661 2046 2358Efficiency (%) 65.8% 65.3% 64.9% 65.3% 59.4% 59.4% 59.3 60.7% 59.6% 60.9% 59.7% 61.1%

Energy balances (MW) Case study


Biomass 225 217 225 225 1096 1098 1088 1089 4383 4339 4350 4345Coal 554 692 562 555 3325 3199 3417 3305 13,054 13,881 13,361 13,388Natural gas 554 357 541 550 1979 2133 1837 1979 8210 7082 7768 7694Butane 12 12 14 14 67 86 75 82 298 246 263 301Gasoline 482 482 482 482 2412 2412 2412 2412 9649 9649 9649 9649Diesel 173 173 173 173 863 863 863 863 3454 3454 3454 3454

37

619

60.

hsocbscaint

TClr

LPG 7 9 7 7 48

Electricity 221 194 200 218 510

Efficiency (%) 65.7% 67.1% 64.3% 65.5% 59.3%

ighlighted in the table using boldface. Carbon that is input to theystem via air is neglected due to the low flow rate relative to thether eight points. Over 99% of the input carbon is supplied from theoal, biomass, and natural gas while the balance is supplied by theutane input to the isomerization and alkylation units. The trendseen in feedstock use from 8 are consistently displayed in the inputarbon flow rates in Table 10. That is, for all of the case studies,

majority of the carbon is input from coal and CO2 sequestration
s highly utilized to reduce the GHG emissions. The biomass andatural gas provide roughly equivalent amounts of input carbono the refineries, which combined represent approximately 40% of
able 10arbon balances (in kg/s) for the optimal solutions for the 24 case studies. Carbon is inpu

iquid product, LPG byproduct, vented CO2, or sequestered (Seq.) CO2. The small amount

esults for the complete superstructure and medium sized capacity (M4) are shown in bo

Case study GDK-S1 GDK-S2 GDK-S3 GDK-S4 GDK-M1 GDK-

Biomass 5.90 5.87 5.88 5.96 29.91 30.58Coal 16.91 17.09 17.17 15.65 89.13 88.96Natural gas 7.32 7.12 7.02 8.47 31.47 30.91Butane 0.19 0.21 0.27 0.19 0.98 0.84Gasoline 7.78 7.78 7.78 7.78 38.91 38.91Diesel 2.85 2.85 2.85 2.85 14.24 14.24Kerosene 1.40 1.40 1.40 1.40 6.98 6.98LPG 0.10 0.08 0.11 0.08 0.46 0.38Vented CO2 0.00 0.00 0.00 0.00 1.15 2.25Seq. CO2 18.20 18.17 18.20 18.16 89.74 88.51% Conversion 40.0% 40.0% 40.0% 40.0% 40.0% 40.0%

Case study GD-S1 GD-S2 GD-S3 GD-S4 GD-M1 GD-M

Biomass 5.99 5.77 5.97 5.99 29.14 29.20Coal 14.54 18.18 14.76 14.58 87.28 83.99Natural gas 9.45 6.09 9.22 9.38 33.75 36.38Butane 0.18 0.18 0.21 0.22 1.01 1.29Gasoline 8.77 8.77 8.77 8.77 43.85 43.85Diesel 3.21 3.21 3.21 3.21 16.04 16.04LPG 0.09 0.11 0.09 0.09 0.58 0.45Vented CO2 0.00 0.00 0.00 0.00 0.00 0.00Seq. CO2 18.10 18.13 18.10 18.10 90.71 90.52% Conversion 40.0% 40.0% 40.0% 40.0% 40.0% 40.0%

45 31 159 176 129 126482 603 1916 1908 1850 2394

3% 59.3 60.6% 58.5% 59.4% 58.6% 60.7%

the input carbon. The output amount of carbon in the total productis constant for each plant capacity, which is consistent with theconstant production capacity that is required for each feedstock-conversion rate. The amount of carbon leaving as LPG is around 1%of that leaving as gasoline, kerosene, and diesel. For all of the casestudies, most of the CO2 generated from the process is capturedand sequestered, with little or no CO2 venting.

For each of the case studies, the carbon conversion rate was set
as a lower bound (i.e., 40%) for the mathematical model. Thus, theconversion of carbon in the four feedstocks to any of the four liquidproducts must be at least as large as the set conversion rate. All of
t to the process via coal, biomass, natural gas, or butanes and exits the process asof CO2 input to the system in the purified oxygen stream (<0.01%) is neglected. Theldface.

M2 GDK-M3 GDK-M4 GDK-L1 GDK-L2 GDK-L3 GDK-L4

29.68 29.39 117.25 118.78 117.00 116.39 81.98 84.91 344.78 330.49 357.45 361.63 38.88 36.06 139.13 154.47 128.30 123.65 1.15 1.11 4.83 3.39 4.57 4.92 38.91 38.91 155.64 155.64 155.64 155.64 14.24 14.24 56.95 56.95 56.95 56.95 6.98 6.98 27.93 27.93 27.93 27.93 0.55 0.46 1.87 2.33 2.41 2.11 0.00 0.00 0.00 0.00 0.00 0.00 91.02 90.88 363.60 364.28 364.39 363.95

40.0 40.0% 40.0% 40.0% 40.0% 40.0%

2 GD-M3 GD-M4 GD-L1 GD-L2 GD-L3 GD-L4

28.94 28.94 116.52 115.37 115.65 115.52 89.70 86.75 342.70 364.41 350.75 351.47 31.33 33.75 140.03 120.79 132.49 131.23 1.13 1.23 4.48 3.70 3.96 4.52 43.85 43.85 175.39 175.39 175.39 175.39

16.04 16.04 64.18 64.18 64.18 64.18 0.55 0.38 1.93 2.15 1.58 1.54

0.00 0.00 0.00 0.00 0.00 0.00 90.66 90.40 362.24 362.57 361.71 361.65

40.0 40.0% 40.0% 40.0% 40.0% 40.0%


Table 11Greenhouse gas (GHG) balances for the optimal solutions for the 24 case studies. The total GHG emissions (in CO2 equivalents – kg CO2eq/s) for feedstock acquisition andtransportation, product transportation and use, CO2 sequestration, and process venting are shown for each study. Process feedstocks include biomass, coal, natural gas, andbutane while products include gasoline, diesel, kerosene, and LPG. The results for the complete superstructure and medium sized capacity (M4) are shown in boldface.

Case study GDK-S1 GDK-S2 GDK-S3 GDK-S4 GDK-M1 GDK-M2 GDK-M3 GDK-M4 GDK-L1 GDK-L2 GDK-L3 GDK-L4

Biomass −27.76 −27.60 −27.68 −28.04 −140.73 −143.88 −139.66 −138.32 −551.76 −558.93 −550.58 −547.68Coal 2.11 2.13 2.14 1.95 11.10 11.08 10.21 10.58 42.94 41.16 44.52 45.04Natural gas 3.42 3.32 3.27 3.95 14.67 14.41 18.13 16.81 64.87 72.02 59.82 57.65Butane 0.02 0.03 0.03 0.02 0.12 0.10 0.14 0.13 0.58 0.41 0.55 0.60Gasoline 30.73 30.73 30.73 30.73 153.64 153.64 153.64 153.64 614.54 614.54 614.54 614.54Diesel 11.17 11.17 11.17 11.17 55.86 55.86 55.86 55.86 223.45 223.45 223.45 223.45Kerosene 5.49 5.49 5.49 5.49 27.46 27.46 27.46 27.46 109.83 109.83 109.83 109.83LPG 0.43 0.35 0.45 0.34 1.89 1.55 2.23 1.87 7.63 9.48 9.80 8.60Vented CO2 0.00 0.00 0.00 0.00 4.23 8.25 0.00 0.00 0.00 0.00 0.00 0.00Seq. CO2 3.33 3.33 3.34 3.33 16.44 16.22 16.68 16.65 66.62 66.74 66.77 66.68Total GHG (kg/bbl) 250.00 250.00 250.00 250.00 250.00 250.00 250.00 250.00 250.00 250.00 250.00 250.00

Case study GD-S1 GD-S2 GD-S3 GD-S4 GD-M1 GD-M2 GD-M3 GD-M4 GD-L1 GD-L2 GD-L3 GD-L4

Biomass −28.18 −27.16 −28.11 −28.18 −137.11 −137.40 −136.17 −136.19 −548.33 −542.91 −547.68 −543.61Coal 1.81 2.26 1.84 1.82 10.87 10.46 11.17 10.80 42.68 45.38 45.04 43.77Natural gas 4.41 2.84 4.30 4.37 15.74 16.96 14.61 15.74 65.29 56.32 57.65 61.19Butane 0.02 0.02 0.03 0.03 0.12 0.16 0.14 0.15 0.54 0.45 0.55 0.55Gasoline 34.62 34.62 34.62 34.62 173.12 173.12 173.12 173.12 692.49 692.49 614.54 692.49Diesel 12.59 12.59 12.59 12.59 62.95 62.95 62.95 62.95 251.79 251.79 223.45 251.79LPG 0.35 0.44 0.36 0.37 2.37 1.85 2.25 1.55 7.87 8.76 109.83 6.27

016

250

twrmbewbMwtc

ieaowttangtGpc(i(uca

elaol

Vented CO2 0.00 0.00 0.00 0.00 0.00

Seq. CO2 3.32 3.32 3.32 3.32 16.62

Total GHG (kg/bbl) 250.00 250.00 250.00 250.00 250.00

he 24 case studies reached this bound, implying that this constraintas active in the optimal solution. Note that this constraint can be

elaxed if a smaller conversion rate of liquid fuels is desired. Ulti-ately, this will have the effect of decreasing the overall fuels cost

y potentially generating additional byproduct electricity. How-ver, recent studies have suggested that the CBGTL process designsill tend to convert between 34% and 37% of the feedstock car-

on when a lower conversion threshold of 25% is set (Baliban, Elia,isener, et al., 2012). Therefore, the minimum threshold of 40%ill serve to provide a baseline measure of comparison between

he case studies while not dramatically impacting the final overallost.

The greenhouse gas (GHG) emission balances for the case stud-es are shown in Table 11. For each of the studies, the total GHGmission target was set to be equal to 50% of the emissions from

standard petroleum based process. For a typical emission levelf 500 kg of CO2 equivalent per barrel, this implies that the totalell-to-wheel GHG emissions for the CBGTL refinery must be less

han 250 kg CO2eq/bbl. The GHG emission rates (in kg CO2eq/s) forhe ten major point sources in the refinery are listed in Table 11nd include (a) acquisition and transportation of the biomass, coal,atural gas, and butane feeds, (b) transportation and use of theasoline, diesel, kerosene, and LPG, (c) transportation and seques-ration of any CO2, and (d) venting of any process emissions. TheHG emissions for feedstock acquisition and transportation in (a),roduct transportation in (b), and CO2 transportation in (c) arealculated from the GREET model for well-to-wheel emissionsArgonne National Laboratory. GREET 1.8b, 2007) and assum-ng transportation distances for feedstocks (50 miles), products100 miles), and CO2 (50 miles). The GHG emissions from prod-ct use in (b) are calculated assuming that each product will beompletly combustion to generate CO2 that is simply vented to thetmosphere.

From Table 11, it is clear that a major component of the lifecyclemissions are attributed to the liquid fuels. In fact, over 80% of the
iquid fuel emissions result from combustion of these fuels in lightnd heavy duty vehicles. The total emissions from transportationf the feedstocks, products, and CO2 represents the balanced of theifecycle emissions for the process. To balance the GHG lifecycle, the
.00 0.00 0.00 0.00 0.00 6.27 0.00

.58 16.61 16.56 66.37 66.43 0.00 66.26

.00 250.00 250.00 250.00 250.00 66.68 250.00

CO2 removed from the atmosphere due to storage in the biomassor storage in the soil is included in the total emissions for biomass.Note that while the net emissions for biomass is negative, therewill still be a positive component to the emissions for biomass har-vesting and transportation. It is important to observe that thoughthe coal was the highest energy input to the refinery, the emissionscontribution from natural gas is higher from coal or biomass.

4. Conclusions

This study has detailed the development of a framework forthe process synthesis of a thermochemical hybrid coal, biomass,and natural gas to liquids plant that incorporates multiple possi-bilities for hydrocarbon production and hydrocarbon upgrading.The framework also included a simultaneous heat, power, andwater integration to compare the costs of utility generation andwastewater treatment in the overall cost of liquid fuels. Thiswork expands previous studies on the CBGTL process (Baliban,Elia, & Floudas, 2012; Baliban et al., 2011) by directly quantify-ing the economic and environmental benefits that are associatedwith (i) Fischer–Tropsch synthesis and subsequent hydrocarbonupgrading and (ii) methanol synthesis, conversion to hydrocar-bons, and subsequent upgrading. The proposed optimization modelwas tested using 24 distinct case studies that are derived fromtwo combinations of products, three plant capacities, and foursuperstructure possibilities. The overall conversion of carbon fromfeedstock to liquid products was selected to be 40% and the green-house gas reduction target was equal to 50% of current petroleumbased refineries. Each case study was globally optimized using abranch-and-bound global optimization algorithm to theoreticallyguarentee that the cost associated with the optimal design waswithin 3–6% of the best value possible.

When producing gasoline, diesel, and kerosene in ratios com-mensurate with Untied States demands, the overall cost ofliquid fuels production ranges from $86/bbl to $94/bbl for small
plants (10,000 barrels per day; kBD), $79/bbl–$88/bbl for mediumplants (50 kBD), and $72/bbl–$80/bbl for large plants (200 kBD).When only gasoline and diesel are produced in a ratio consis-tent with national demand, the cost decreases for each of the

4 Chem

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A

N

Aw

pglpFrF2

A

damotid

u

P

i

s

structure}


apacities to a range of $85/bbl–$93/bbl for small, $78/bbl–$86/bblor medium, and $71/bbl–$78/bbl for large plants. This decreasen cost is generally due to the reduction in investment neededo fractionate and convert the distillate to diesel only opposedo both diesel and kerosene. For the four different superstructureossibilities investigated in this study, it is evident that FT syn-hesis followed by fractionation (superstructure 1) and upgradings more expensive than FT synthesis followed by catalytic ZSM-

conversion to gasoline-range hydrocarbons (superstructure 2).dditionally, methanol synthesis, conversion to hydrocarbons, andubsequent upgrading (superstructure 3) is consistently cheaperhan FT synthesis for all capacity levels. This is due to the decreasen investment cost associated with hydrocarbon production andpgrading when compared to FT synthesis. These findings indicatehat the methanol route is preferential to the FT route when follow-ng an “either or” logic. However, investigation of a “combination”uperstructure that considered all of the topologies (superstruc-ure 4) in superstructures 1–3 indicates that a combination of FTynthesis and methanol synthesis will provide the lowest overallost. In this case, the MTG route provides a majority of the gasolinehile a majority of the distillate (diesel and kerosene) is gener-

ted through fractionation and refining of the FT effluent. Thoughver 80% of the final hydrocarbons were produced via the methanolynthesis route, the final process topologies show that the abilityo consume CO2 in iron-based FT reactors helps to reduce feedstockosts and therefore provide an economic advantage over a topologyhat utilizes only methanol synthesis.

cknowledgements

The authors acknowledge partial financial support from theational Science Foundation (NSF EFRI-0937706).

ppendix A. Mathematical model for process synthesisith simultaneous heat, power, and water integration

The nomenclature for all terms in the mathematical model forrocess synthesis with simultaneous heat, power, and water inte-ration is shown below. All constraints included in the model areisted subsequently with a corresponding description of how thatarticular equation governs proper operation of the process design.or a more extensive discussion of the mathematical model, theeader is directed to previously published works (Baliban, Elia, &loudas, 2012; Baliban, Elia, Misener, et al., 2012; Baliban et al.,011).

.1. Process units

The set of units, U, is presented in full detail in Table A1 andefined formally in Eq. (A.1). Note that several units in Table A1re listed as un. The n subscript represents the consideration ofultiple forms of the same process unit, each with a distinct set

f operating conditions (e.g., temperature and pressure). Thoughhese unit properties are generally given as continuous variablesn a process synthesis problem, they have been assumed to takeiscrete choices and will be modeled using binary variables.

∈ U = {Complete set of process units listed in Table A1} (A.1)

rocess species

The set of all species, S, is listed in Table A2 and defined formallyn Eq. (A.2).

∈ S = {Complete set of species listed in Table A2} (A.2)


Indices/sets

The indices are used throughout the mathematical model arelisted below.

u : Process unit indexs : Species indexa : Atom indexp : Proximate analysis indexr : Reaction indexi : General counting index

The set, U, is defined as the complete set of process units. Severalsubsets of units are then defined for specific areas of the CBGTLprocess as presented below.

uBGS = {u : u = BGSn}uCGS = {u : u = CGSn}uRGS = {u : u = RGSn}uATR = {u : u = ATRn}

The set of all atoms, A, includes C, H, O, N, S, Cl, Ar, and a genericAsh atom. Typically, the biomass and coal ash will consist of mul-tiple metal oxides, but the ash is assumed to be inert in the CBGTLprocess, so the treatment of the ash as an atomic element is justi-fied.

a ∈ A = {C, H, O, N, S, Cl, Ar, Ash}

The list of all unit connections, UC, is derived below.

UC = {(u, u′) : ∃ a connection between unit u and unit u′ in the super

Using a priori knowledge about the operations of each unit inthe CBGTL process, the complete set of species that can possiblyexist in a stream from unit u to unit u′ is defined as SUC

u,u′ . The set (u,

u′, s) ∈ SUF is then constructed from all streams in UC along with theset of all species s that exist within a given unit u (SU).

SUF = {(u, u′, s) : ∃s ∈ SUCu,u′ }

SU = {(s, u) : ∃(u, u′, s) ∈ SUF or ∃(u′, u, s) ∈ SUF }

Parameters

With the exception of all biomass and coal species, char, and thepseudocomponents, the molecular formula is equal to the speciesindex defined in Table A2. The pseudocomponent hydrocarbonsand oxygenate formulas are given by Bechtel while the formulas forbiomass and coal compounds are derived from the ultimate analy-sis and normalized to one mole of carbon. Char has been assumedto consist completely of carbon and ash has been assigned a genericmolecular weight of 1.0 g/mol. The atomic ratio (ARs,a) of atom a inspecies s is derived from the molecular formulas in Table A2.

ARs,a : Atomic ratio of atom a in species s

Using the appropriate atomic weight of atom a (AWa), the molec-ular weight of all species s (MWs) is defined using Eq. (A.3).

AWa : Atomic weight of atom a

MW =∑

AW · AR(A.3)

s

a

a s,a

The proximate analysis for the biomass and coal species s isdescribed by the total mass of moisture per unit mass of dry input


Table A1Process units present in the CBGTL synthesis problem. The subscript n corresponds to multiple forms of the same process unit, each with a distinct set of operating conditionsor ratios of feedstock. Distinct process units are used in lieu of continuous variables representing the process operating conditions. This will prevent the use of bilinear termswhen specifying feedstock ratios or highly non-linear equations when specifying equilibrium constants or species enthalpies.

Unit name Unit index Unit name Unit index

Process inletsInlet coal INCOAL Inlet Natural gas INNG

Inlet biomass INBIO Inlet air INAIR

Inlet water INH2O Inlet butane INBUT

Process outletsOutlet gasoline OUTGAS Outlet diesel OUTDIE

Outlet kerosene OUTKER Outlet ash OUTASH

Outlet sulfur OUTS Outlet scrubbed HCl OUTSCR

Outlet vent OUTV Outlet propane OUTPRO

Outlet sequestered CO2 OUTCO2 Outlet Wastewater OUTWW

Syngas generationBiomass dryer BDR Biomass dryer air heater XBDR

Biomass lockhopper BLK Biomass Gasifier BGSn

First biomass vapor cyclone BC1 Second biomass vapor cyclone BC2

Tar cracker TCK Tar cracker splitter SPTCK

Tar cracker cooler XTCK Coal dryer CDRCoal dryer air heater XCDR Coal lockhopper CLKCoal gasifier CGSn First coal vapor cyclone CC1

Second coal vapor cyclone CC2 Second coal cyclone splitter SPCC2

Second coal cyclone cooler XTCK

Syngas cleaningReverse water gas shift unit RGSn RGS effluent cooler XRGS

COS–HCN hydrolyzer CHH HCl scrubber HSCAcid gas flash vapor cooler XAGF Acid gas flash 2-phase cooler XAGFn

Acid gas flash unit AGF Acid gas thermal analyzer XAGR

Acid gas removal unit AGR First CO2 compressor CO2CCO2 recycle compressor CO2RC CO2 sequestration compressor CO2SCAcid gas compressor AGC

Claus sulfur recoveryAcid gas splitter SPAG Acid gas preheater XAG

Claus combustor CC First sulfur converter SC1

First sulfur separator SS1 Second sulfur converter heater XSC2

Second sulfur converter SC2 Second sulfur separator SS2

Third sulfur converter heater XSC3 Third sulfur converter SC3

Third sulfur separator SS3 Sulfur pit SPTTail gas hydrolyzer TGH Tail gas flash vapor cooler XTGF

Tail gas flash 2-phase cooler XTGFn Tail gas flash unit TGFTail gas compressor TGC

Hydrocarbon productionMTFTWGS-N Iron MT fWGS nominal wax FT MTFTWGS-M Iron MT fWGS minimal wax FTFT-ZSM5 ZSM-5 hydrocarbon conversion unit ZSM5F ZSM-5 product fractionationMEOHS Methanol synthesis unit MEOH-F Methanol flash unitMEDEG Methanol degasser MTG Methanol to gasoline ZSM-5 reactorMTO Methanol to olefins ZSM-5 reactor MTO-F MTO fractionationOGD Olefins to gasoline/distillate MTODF OGD fractionationFischer–Tropsch compressor FTC Fischer–Tropsch splitter SPFT

Low-temperature preheater XLTFT Low-temperature splitter SPLTFT

Low-temperature iron-based FT LTFT Low-temperature cobalt-based FT LTFTRGSHigh-temperature preheater XHTFT High-temperature splitter SPHTFT

High-temperature iron-based FT HTFT High-temperature cobalt-based FT HTFTRGSLow-temperature effluent cooler XLTFTC High-temperature effluent cooler XHTFTC

Water-soluble oxygenates separator WSOS Vapor-phase oxygenates separator VPOSPrimary vapor–liquid–water separator VLWS

Hydrocarbon recoveryHydrocarbon recovery column HRC Wax Hydrocracker WHCDistillate hydrotreater DHT Kerosene hydrotreater KHTNaphtha hydrotreater NHT Naphtha reformer NRFC4 Isomerizer C4I C5–C6 Isomerizer C56IC3–C4–C5 Alkylation unit C345A Saturated gas plant SGPDiesel blender DBL Gasoline blender GBLHCKO1 Mixed hydrocarbon knockout 1 HCKO2 Mixed hydrocarbon knockout 2DEETH De-ethanizer ABS-COL Absorber columnCO2SEP 1-stage Rectisol CO2 separation STA-COL Stabilizer columnALK-UN HF alkylation unit LPG-ALK LPG/Alkylate splitterSP-COL Splitter column

Recycle gas treatmentLight gas compressor LGC Light gas splitter SPLG

Auto-thermal reactor ATRn Auto-thermal reactor splitter SPATRn

Fuel combustor FCM Fuel combuster effluent cooler XFCM

Fuel combustor flash unit FCF First gas turbine air compressor GTAC1


Table A1 (Continued)

Unit name Unit index Unit name Unit index

Second gas turbine air compressor GTAC2 Gas turbine combustor GTCFirst gas turbine GT1 Second gas turbine GT2

Gas turbine effluent cooler XGT Gas turbine flash unit GTFGas turbine effluent compressor GTEC CO2 recovery unit CO2RWater gas shift unit WGS

Water treatmentSour stripper SS Sour gas compressor SGCBiological digestor BD Reverse osmosis ROCooling tower CLTR Process cooling COOL-PHeat &Power system HEP Heat &Power utilities HEAT-PDeaerator DEA Process water economizer XWPR

Process water boiler XWBL

Hydrogen/oxygen productionPSA effluent splitter SPPSA Pressure-swing Absorption unit PSAPSA hydrogen preheater XH2P PSA hydrogen splitter SPH2P

Electrolyzer EYZ Electrolyzer oxygen preheater XO2E

Electrolyzer oxygen splitter SPO2E Electrolyzer hydrogen preheater XH2E

Electrolyzer hydrogen splitter SPH2E Air Compressor AC

(a

a

TSb

Air separation unit ASUASU oxygen preheater XO2A

OC oxygen preheater XO2C

PAMs ) and the dry weight fractions (PAD

p,s) of the ash, fixed carbon,nd volatile matter components p.

PAM : Mass of water per unit mass of dry species s
s
PADp,s : Dry mass fraction of proximate analysis component p in species s

In this study, switchgrass was chosen for the biomass feedstocknd low-volatile bituminous coal was chosen for the coal feedstock.

able A2pecies present in the CBGTL synthesis problem. The molecular formula of the pseudocoiomass and coal species are derived from the ultimate analysis assuming that the “atom

Species name Species index Species name

Acid gasesSulfur dioxide SO2 Hydrogen sulfur

Hydrogen cyanide HCN Ammonia

Carbon dioxide CO2

Light non-hydrocarbon gasesOxygen O2 Nitrogen

Nitric oxide NO Nitrous oxide

Carbon monoxide CO Hydrogen

HydrocarbonsMethane CH4 Acetylene

Ethane C2H6 Propylene

Isobutylene iC4H8 1-Butene

n-Butane nC4H10 1-Pentene

n-Pentane nC5H12 1-Hexene

n-Hexane nC6H14 1-Heptene

1-Octene C8H16 n-Octane

n-Nonane C9H20 1-Decene

1-Undecene C11H22 n-Undecane

n-Dodecane C12H26 1-Tridecene

1-Tetradecene C14H28 n-Tetradecane

n-Pentadecane C15H32 1-Hexadecene

1-Heptadecene C17H34 n-Heptadecane

n-Octadecane C18H38 1-Nonadecene

1-Eicosene C20H40 n-Eicosane

C22 Pseudocomponent C22OP C23 Pseudocomponent



VP Oxygenate OXVAP HP Oxygenate

ProductsGasoline GAS Diesel

Solid sulfur S

Non-conventional componentsBiomass e.g. Perennial Coal

Feedstock ash Ash

Oxygen compressor OCOC Oxygen splitter SPO2C

Variables

Continuous variables are used in the mathematical model toS
describe the species molar flow rates (Nu,u′,s), the total molar flow
rates (NTu,u′ ), the extent of reaction in a process unit (�u

r ), the molar

composition of a stream (xSu,u′,s), the split fraction of a stream

between two units (spu,u′ ), the total stream enthalpy flow rate

mponent hydrocarbons and oxygenates are given by Bechtel. The formula for theic” weight of ash is 1.0 g/mol.

Species index Species name Species index

H2S Carbonyl sulfide COS

NH3 Hydrogen chloride HCl

N2 Argon ArN2O Water H2OH2

C2H2 Ethylene C2H4

C3H6 Propane C3H8

nC4H8 Isobutane iC4H10

C5H10 2-Methylbutane iC5H12

C6H12 2-Methylpentane iC6H14

C7H14 n-Heptane C7H16

C8H18 1-Nonene C9H18

C10H20 n-Decane C10H22

C11H24 1-Dodecene C12H24

C13H26 n-Tridecane C13H28

C14H30 1-Pentadecene C15H30

C16H32 n-Hexadecane C16H34

C17H36 1-Octadecene C18H36

C19H38 n-Nonadecane C19H40

C20H42 C21 Pseudocomponent C21OPC23OP C24 Pseudocomponent C24OPC26OP C27 Pseudocomponent C27OPC29OP C30+ Pseudocomponent C30WaxOXHC AP Oxygenate OXH2O

DIE Kerosene KER

e.g. LV-bituminous Gasifier char Char

Chem

(

atasc

ocatF

y

G

M

(

�

(

N

P

N

(


HTu,u′ ), the heat lost from a unit (Q L

u ), the heat transferred to or

bsorbed from a unit (Qu), the delivered cost of feedstock (CostFs ),

he cost of CO2 sequestration (CostSeq), the cost of electricity (CostEl),nd the levelized unit investment cost (CostU

u ). Note that the sub-cripts u and u′ are both used to denote an element of the set U andan be used interchangeably in the stream flow indices.

NSu,u′,s : Molar flow of species s from unit u to unit u′

NTu,u′ : Total molar flow from unit u to unit u′

�ur : Extent of reaction r in unit u

xSu,u′,s : Molar composition of species s from unit u to unit u′

spu,u′ : Split fraction of stream going from unit u to unit u′

HTu,u′ : Total enthalpy flow from unit u to unit u′

Q Lu : Heat lost from unit u

Qu : Heat transferred to or absorbed from unit u

CostFs : Total delivered cost of feedstock s

CostSeq : Total sequestration cost of CO2

CostEl : Total cost of electricityCostU

u : Total levelized cost of unit u

Binary variables (yu) are introduced to represent the logical usef a process unit u. These binary variables are only needed for spe-ific process units since many of the units in the CBGTL process willlways be required. The units that require binary variables includehe biomass and coal gasifiers, the reverse water gas shift unit, theischer–Tropsch units, the autothermal reactor, and the gas turbine.

u : Logical existence of process unit u (i.e., it takes the valueof one if unit u is selected and zero otherwise)

eneral constraints

ass balancesSpecies balances

∑u′,u)∈UC

NSu′,u,s −

∑(u,r,s′)∈RU

�r,s

�r,s′· �u

r −∑

(u,u′)∈UC

NSu,u′,s

= 0 ∀s ∈ SUu , u ∈ UBal

Sp (A.4)

Extent of reaction

ur − fcu

r ·∑

(u′,u,s)∈SUF

NSu′,u,s = 0 ∀(u, r, s) ∈ RU (A.5)

Atom balances∑u′,u,s)∈SUF

ARs,a · NSu′,u,s −

∑(u,u′,s)∈SUF

ARs,a · NSu,u′,s

= 0 ∀a ∈ AUu , u ∈ UBal

At (A.6)

Total mole balance

Tu′,u −

∑(u,u′,s)∈SUF

NSu′,u,s = 0 ∀(u, u′) ∈ UC (A.7)

rocess splittersSet unit split fractions

Su,u′,s − xS

uI ,u,s · NTu,u′ = 0 ∀(u, u′, s) ∈ SUF , u ∈ USp (A.8)

Split fractions sum to 1∑u,u′,s)∈SUF

xSu,u′,s − 1 = 0 ∀(u, u′) ∈ UCComp (A.9)


Flash unitsUpper bound of liquid phase split fraction

xSu,uL,s − min{1,

1

KVLEu,s

} ≤ 0 ∀(u, uL, s) ∈ SUF , u ∈ UFl (A.10)

Upper bound of vapor phase split fraction

xSu,uV ,s − min{1, KVLE

u,s } ≤ 0 ∀(u, uV , s) ∈ SUF , u ∈ UFl (A.11)

Set liquid phase split fraction

xSu,uL,s · NT

u,uL− NS

u,uL,s = 0 ∀u ∈ UFl (A.12)

Set vapor phase split fraction

xSu,uV ,s · NT

u,uV− NS

u,uV ,s = 0 ∀u ∈ UFl (A.13)

Set phase equilibrium

xSu,uV ,s − KVLE

u,s · xSu,uL,s = 0 ∀u ∈ UFl (A.14)

Heat balancesConservation of energy

∑(u,u′)∈UC

HTu,u′ −

∑(u′,u)∈UC

HTu′,u − Qu − Q L

u − Wu = 0 ∀u ∈ U

UAgg(A.15)

Total heat balance

HTu,u′ −

∑(u,u′,s)∈SUF

HSu,u′,s = 0 ∀(u, u′) ∈ UC (A.16)

Logical unit existenceBound on molar flows

∑(u′,u)∈UC

NTu′,u − UBN

u · yu ≤ 0 ∀u ∈ UEx (A.17)

Upper bound on inlet enthalpy flow

HTu′,u − UBH

u′,u · yu ≤ 0 ∀(u′, u) ∈ UC, u ∈ UEx (A.18)

Lower bound on inlet enthalpy flow

LBHu′,u · yu − HT

u′,u ≤ 0 ∀(u′, u) ∈ UC, u ∈ UEx (A.19)

Upper bound on outlet enthalpy flow

HTu,u′ − UBH

u′,u · yu ≤ 0 ∀(u, u′) ∈ UC, u ∈ UEx (A.20)

Lower bound on outlet enthalpy flow

LBHu,u′ · yu − HT

u,u′ ≤ 0 ∀(u, u′) ∈ UC, u ∈ UEx (A.21)

Process inlets

Feedstock moisture contentSet biomass moisture content from proximate analysis

MSu,u′,H2O −

∑s∈SBio

PAMs · MS

u,u′,s = 0 (u, u′) = (INBIO, BDR) (A.22)

Set coal moisture content from proximate analysis

MSu,u′,H2O −

∑s∈SCoal

PAMs · MS

u,u′,s = 0 (u, u′) = (INCOAL, CDR) (A.23)

Known stream compositionsSet stream compositions for inlet streams

NSu,u′,s − xK

u,s · NTu,u′ = 0 ∀(u, u′, s) ∈ SUF , u = {INAIR, INNG, INBUT}

(A.24)

5 Chem

C

r

s

G

G

G

G

G

G

P

M

M

S

B

M

M

M


oal to natural gas ratioSet coal to natural gas inlet ratio based on lower heating value

atios

∑∈SCoal

NSINCOAL,CDR,s · LHVs − LHVRat

CG · LHVNG ·∑

(INNG,u)∈UC

NTINNG,u = 0

(A.25)

reenhouse gas emissions reductionSet reduction from petroleum based processes

HGCBGTL − GHGRed · GHGPet = 0 (A.26)

Sum emissions from CBGTL components

HGCBGTL − GHGSeq − GHGProc − GHGFeed = 0 (A.27)

Set emissions from feedstock acquisition

HGFeed −∑u∈UIn

∑(u,u′,s)∈SUF

GHGTs · MS

u,u′,s = 0 (A.28)

Set emissions from CO 2 sequestration

HGSeq − GHGTCO2

· MWCO2 · NSCO2SC,OUTCO2

,CO2= 0 (A.29)

Set emissions from CO2 venting

HGProc − MWCO2 · NSCO2R,OUTV,CO2

= 0 (A.30)

rocess outlet fuel ratios

Set gasoline to diesel output ratio

WGAS · NSGBL,OUTGAS,GAS − RatG−D · MWDIE · NS

DBL,OUTDIE,DIE = 0

(A.31)

Set diesel to kerosene output ratio

WDIE · NSDBL,OUTDIE,DIE − RatD−K · MWKER · NS

KHT,OUTKER,KER = 0

(A.32)

yngas generation

iomass/coal driersUpper bound for biomass drier activation

Su,u′,H2O − MTBio · MT

u,u′ − UB · yu ≤ 0 (u, u′) = (INBIO, BDR)

(A.33)

Upper bound for coal drier activation

Su,u′,H2O − MTCoal · MT

u,u′ − UB · yu ≤ 0 (u, u′) = (INCOAL, CDR)

(A.34)

Lower bound for biomass drier activation

TBio · MTu,u′ − MS

u,u′,H2O − UB · (1 − yu) ≤ 0 (u, u′) = (INBIO, BDR)

(A.35)


Lower bound for coal drier activation

MTBio · MTu,u′ − MS

u,u′,H2O − UB · (1 − yu) ≤ 0 (u, u′) = (INCOAL, CDR)

(A.36)

Upper bound for biomass drier moisture evaporation

MTBio · MTu,u′ − MS

u,u′,H2O − UB · (1 − yu) ≤ 0 (u, u′) = (BDR, BLK)

(A.37)

Lower bound for biomass drier moisture evaporation

MSu,u′,H2O − MTBio · MT

u,u′ − UB · (1 − yu) ≤ 0 (u, u′) = (BDR, BLK)

(A.38)

Upper bound for coal drier moisture evaporation

MTCoal · MTu,u′ − MS

u,u′,H2O − UB · (1 − yu) ≤ 0 (u, u′) = (CDR, CLK)

(A.39)

Lower bound for coal drier moisture evaporation

MSu,u′,H2O − MTCoal · MT

u,u′ − UB · (1 − yu) ≤ 0 (u, u′) = (CDR, CLK)

(A.40)

Gasifier lockhoppersSet CO2 lockhopper flow rate

MSCO2C2,BLK,CO2

− mf u ·∑s∈SBio

MSBDR,BLK,s = 0 (A.41)

Biomass gasifierWater–gas-shift equilibrium

Nu,BC1,CO · Nu,BC1,H2O − KRGSu · Nu,BC1,CO2 · Nu,BC1,H2 = 0 ∀u ∈ UBGS

(A.42)

Hydrocarbon conversion fraction

Mu,BC1,s −∑

(u′,u,s)∈SUF

cf HCu,s · MS,Calc

s = 0 ∀s ∈ SHC, u ∈ UBGS (A.43)

Hydrocarbon generation from pyrolysis

MS,Calcs −

∑s′∈SBio

∑(u′,u,s′)∈SUF

PyrHCs,s′ · MS

u′,u,s′−∑

(u′,u)∈UC

MSu′,u,s=0 u ∈ UBGS

(A.44)

Set ratio of NO to N2O

Nu,BC1,NO − sru, NON2O

· Nu,BC1,N2O = 0 ∀u ∈ UBGS (A.45)

Set ratio of HCN to NH3

Nu,BC1,HCN − sru, HCNNH3

· Nu,BC1,NH3 = 0 ∀u ∈ UBGS (A.46)

Set amount input nitrogen to NH3 and N2

Nu,BC1,NH3 + 2 · Nu,BC1,N2 − nf u ·∑

(u,BC1,s)∈SUF

NSu,BC1,s · ARs,N

= 0 ∀u ∈ UBGS (A.47)

Chem

N

N

M

N

Q

u

B

r

r

C

M

N

M

M S

)

N


Set ratio of NH3 to N2

u,BC1,NH3 − (a1u,N2

+ a2u,N2

· Tu) · (Nu,BC1,NH3 + 2 · Nu,BC1,N2 )

= 0 ∀u ∈ UBGS (A.48)

Set ratio of COS to H2S

u,BC1,COS − sru, COSH2S

· Nu,BC1,H2S = 0 ∀u ∈ UBGS (A.49)

Amount of char production

WChar · NSu,BC1,Char − (a1

u,Char + a2u,Char · Tu) ·

∑s∈SBio

MWs · NSBLK,u,s

= 0 ∀u ∈ UBGS (A.50)

Rate of ash removal

Su,OUTASH,Ash − sf u,Ash ·

∑(u′,u)∈UC

NSu′,u,Ash = 0 ∀u ∈ UBGS (A.51)

Gasifier heat loss

Lu + hlu ·

∑s∈SBio

MSBLK,u,s · LHVs = 0 ∀u ∈ UBGS (A.52)

Logical use of one gasifier temperature∑∈UBGS

yu − 1 = 0 (A.53)

iomass gasifier solidsRemoval of solids from first cyclone

f BC1 · NTBGS,BC1 − NT

BC1,BGS = 0 (A.54)

Removal of solids from second cyclone

f BC2 · NTBC1,BC2 − NT

BC2,BGS = 0 (A.55)

oal gasifierSet CO2 lockhopper flow rate

SSPCO2,CLK,CO2

− mf u ·∑

s∈SCoal

MSCDR,CLK,s = 0 (A.56)

Water–gas-shift equilibrium

u,CC1,CO · Nu,CC1,H2O − KRGSu · Nu,CC1,CO2 · Nu,CC1,H2 = 0 ∀u ∈ UCGS

(A.57)

Hydrocarbon conversion fraction

u,CC1,s −∑

(u′,u,s)∈SUF

cf HCu,s · MS,Calc

s = 0 ∀s ∈ SHC, u ∈ UCGS (A.58)

Hydrocarbon generation from pyrolysis

S,Calcs −

∑s′∈SCoal

∑(u′,u,s′)∈SUF

PyrHCs,s′ · MS

u′,u,s′−∑

(u′,u)∈UC

MSu′,u,s=0 u ∈ UCG

(A.59

Set ratio of NO to N2O

u,CC1,NO − sr NON2O

· Nu,CC1,N2O = 0 ∀u ∈ UCGS (A.60)


Set ratio of HCN to NH3

Nu,CC1,HCN − sr HCNNH3

· Nu,CC1,NH3 = 0 ∀u ∈ UCGS (A.61)

Set amount input nitrogen to NH3 and N2

Nu,CC1,NH3 + 2 · Nu,CC1,N2 − nf u ·∑

(u,CC1,s)∈SUF

NSu,CC1,s · ARs,N

= 0 ∀u ∈ UCGS (A.62)

Set ratio of NH3 to N2

Nu,CC1,NH3 − (a1u,N2

+ a2u,N2

· Tu) · (Nu,CC1,NH3 + 2 · Nu,CC1,N2 )

= 0 ∀u ∈ UCGS (A.63)

Set ratio of COS to H2S

Nu,CC1,H2S − sru, COSH2S

· Nu,CC1,COS = 0 ∀u ∈ UCGS (A.64)

Amount of char production

MWChar · NSu,CC1,Char − (a1

u,Char + a2u,Char · Tu) ·

∑s∈SCoal

MWs · NSCLK,u,s

= 0 ∀u ∈ UCGS (A.65)

Rate of ash removal

NSu,OUTASH,Ash − sf u,Ash ·

∑(u′,u)∈UC

NSu′,u,Ash = 0 ∀u ∈ UCGS (A.66)

Gasifier heat loss

Q Lu + hlu ·

∑s∈SCoal

MSCLK,u,s · LHVs = 0 ∀u ∈ UCGS (A.67)

Logical use of one gasifier temperature∑u∈UCGS

yu − 1 = 0 (A.68)

Coal gasifier solidsRemoval of solids from first cyclone

rf CC1 · NTCGS,CC1 − NT

CC1,CGS = 0 (A.69)

Removal of solids from second cyclone

rf CC2 · NTCC1,CC2 − NT

CC2,CGS = 0 (A.70)

Syngas cleaning

Reverse water–gas-shift unitBypass of inert species∑

(u′,u,s)∈SUF

NSu′,u,s −

∑(u,u′,s)∈SUF

NSu,u′,s

= 0 ∀s ∈ SInu , u ∈ URGS (A.71)


Nu′,u,CO · Nu′,u,H2O − KRGSu′ · Nu′,u,CO2 · Nu′,u,H2

= 0 ∀u′ ∈ URGS, u = XRGS (A.72)

5 Chem

u

C

(

N

N

A

N

N

C

(

H

F

(

(

W

W

W


Logical use of unit with at most one temperature∑∈URGS

yu − 1 ≤ 0 (A.73)

OS–HCN hydrolyzerBypass of inert species∑

u′,u,s)∈SUF

NSu′,u,s −

∑(u,u′,s)∈SUF

NSu,u′,s = 0 ∀s ∈ SIn

u , u ∈ UCHH (A.74)

COS–H2S equilibrium

u′,u,COS · Nu′,u,H2O − KCOSu′ · Nu′,u,CO2 · Nu′,u,H2S

= 0 (u′, u) = (CHH, HSC) (A.75)

HCN–NH3 equilibrium

u′,u,HCN · Nu′,u,H2O − KHCNu′ · Nu′,u,CO · Nu′,u,NH3

= 0 (u′, u) = (CHH, HSC) (A.76)

cid gas recoverySet CO2 molar fraction in clean output

SAGR,SPAGR,CO2

− rf AGR · NTAGR,SPCG

= 0 (A.77)

Set CO2 output flow rates

TAGR,CO2C − sf AGR · (NT

AGR,CO2C + NTAGR,MXCO2RC

) = 0 (A.78)

laus sulfur recovery

Set inlet combustor oxygen level∑

u,CC)∈UC

NSu,CC,O2

− erCC ·∑

(u,CC,s)∈SUF

NSu,CC,s · sors = 0 (A.79)

ydrocarbon production

ischer–TropschSet ratio of H2 to CO in cobalt-based inlet∑

u′,u,H2)∈SUF

− FTRu,CO ·∑

(u′,u,CO)∈SUF

= 0 ∀u ∈ UCoFT (A.80)

Set ratio of H2 to CO and CO2 in iron-based inlet∑

u′,u,H2)∈SUF

− FTRu,CO ·∑

(u′,u,CO)∈SUF

− FTRu,CO2 ·∑

(u′,u,CO2)∈SUF

= 0 ∀u ∈ UIrFT (A.81)

Adjust weight fraction of C1 species

1 = 12

(1 −∞∑

n=5

Wn) (A.82)


2 = 16

(1 −∞∑

n=5

Wn) (A.83)


3 = 16

(1 −∞∑

n=5

Wn) (A.84)



W4 = 16

(1 −∞∑

n=5

Wn) (A.85)

Set weight fraction of Cn species from Anderson–Schultz–Flory dis-tribution

Wn = n(1 − ˛)2˛n−1 ∀5 ≤ n ≤ 29 (A.86)

Set weight fraction of wax

WWax =∞∑

n=30

n(1 − ˛)2˛n−1 (A.87)

Set carbon distribution from weight fractions

crn = n · Wn∑29n=1n · Wn + nWax · WWax

(A.88)

Set exactly one low-temperature unit

yLTFT + yLTFTRGS − 1 = 0 (A.89)

Set exactly one high-temperature unit

yHTFT + yHTFTRGS − 1 = 0 (A.90)

Aqueous phase oxygenates separatorRemoval of aqueous phase oxygenates

NSWSOS,VLWS,s = 0 ∀s ∈ SAPO (A.91)

Vapor phase oxygenates separatorRemoval of vapor phase oxygenates

NSVPOS,HRC,s = 0 ∀s ∈ SVPO (A.92)

Hydrocarbon upgrading

Hydrocarbon upgrading unitsSet carbon distribution fractions of total input

NSu,u′,s · ARs,C − cf u,u′,s ·

∑(u′′,u,s′)∈SUF

NSu′′,u,s′ · ARs′,C

= 0 ∀u ∈ UUG, (u, u′, s) ∈ sUF (A.93)

Saturated gas plantSet fractional recovery of light gases

NSSGP,C4I,s − rf s ·

∑(u,SGP,s)∈SUF

NSu,SGP,s = 0 ∀s ∈ SC4 (A.94)

Recycle gas treatment

Fuel combustorSet inlet combustor oxygen level∑

(u,FCM)∈UC

NSu,FCM,O2

− erFCM ·∑

(SPLG,FCM,s)∈SUF

NSSPLG,FCM,s · sors = 0

(A.95)

Auto-thermal reactor
Logical use of one temperature
∑u∈UATR

yu − 1 = 0 (A.96)

Chem

x

x

x

x

(

G

N

N

e

Q

W

S

N

N



NSu,u′,CO2

· NSu,u′,H2

− KRGSu · NS

u,u′,CO · NSu,u′,H2O

= 0 ∀(u, u′) ∈ UC, u ∈ UATR (A.97)

CH4 Steam reforming equilibrium

Su,u′,CO · xS

u,u′,H2

3 −KSRu,CH4

· xSu,u′,CH4

· xSu,u′,H2O

= 0 ∀(u, u′) ∈ UC, u ∈ UATR (A.98)

C2H2 steam reforming equilibrium

Su,u′,C2H2

· xSu,u′,H2O −

KSRu,CH4

KSRu,C2H2

· xSu,u′,CH4

· xSu,u′,CO

= 0 ∀(u, u′) ∈ UC, u ∈ UATR (A.99)

C2H4 steam reforming equilibrium

Su,u′,C2H4

−KSR

u,C2H2

KSRu,C2H4

· xSu,u′,C2H2

· xSu,u′,H2

= 0 ∀(u, u′) ∈ UC, u ∈ UATR

(A.100)

C2H6 Steam reforming equilibrium

Su,u′,C2H6

−KSR

u,C2H4

KSRu,C2H6

· xSu,u′,C2H4

· xSu,u′,H2

= 0 ∀(u, u′) ∈ UC, u ∈ UATR

(A.101)

Bypass of inert species∑

u′,u,s)∈SUF

NSu′,u,s −

∑(u,u′,s)∈SUF

NSu,u′,s = 0 ∀u ∈ UATR, s ∈ SIn

ATR (A.102)

as turbineSet air leakage from first compressor

SGTAC1,OUTV,s − lkGTAC1

· NSINAIR,GTAC1,s = 0 ∀(GTAC1, s) ∈ SU

(A.103)

Set air bypass from first compressor

SGTAC1,GT2,s − byGTAC1

· NSINAIR,GTAC1,s = 0 ∀(GTAC1, s) ∈ SU

(A.104)

Set inlet oxygen flow rate in combustor

rGTC ·∑

(u,GTC,s)∈SUF

sors · NSu,GTC,s −

∑(u,GTC,s)∈SUF

NSu,GTC,O2

= 0 (A.105)

Set heat loss in combustor

LGTC − hlGTC · (HT

SPLG,GTC − HTXGTF,GTF) = 0 (A.106)

astewater treatment

our stripperSet recovery fraction of H2O in bottoms

SSS,SPSS,H2O − rf SS,H2O ·

∑NS

u,SS,s = 0 (A.107)

(u,SS)∈UC

Set fraction of sour species in bottoms

SSS,SPSS,s − xKn

SS,SPSS,s · NTSS,SPSS,s = 0 ∀(SS, SPSS, s) ∈ SUF (A.108)


Energy balance using reboiler and condensor

Q RebSS + Q Cond

SS − QSS = 0 (A.109)

Set energy use for reboiler and condensor

HRSS · Q RebSS + Q Cond

SS = 0 (A.110)

Biological digestorSet biogas ratio of CH4 to CO2

NSBD,CC,CH4

− crBD · NSBD,CC,CO2

= 0 (A.111)

Reverse osmosisSet removal fraction of solids

NSRO,SPRO,s − rf RO · NS

MXRO,RO,s = 0 ∀s ∈ SSol (A.112)

Cooling cycleCooling tower flow rate from energy requirement

QC − hrCOOL−P · NSCLTR,COOL−P,H2O = 0 (A.113)

Cooling tower evaporation loss

NEvapCLTR − 0.00085 · �TCLTR · NS

CLTR,COOL−P,H2O = 0 (A.114)

Cooling tower drift loss

NDriftCLTR − 0.001 · NS

MXCLTR,CLTR,H2O = 0 (A.115)

Sum total cooling tower losses

NEvapCLTR + NDrift

CLTR − NSCLTR,OUTV,H2O = 0 (A.116)

Set known cooling tower output solid concentrations

xKnCLTR,SPCLTR,s · NT

CLTR,SPCLTR− NS

CLTR,SPCLTR,s = 0 ∀s ∈ SSol (A.117)

Steam cycleSet known process steam boiler output solid concentrations

xKnXPWB,MXBLR,s · NT

XPWB,MXBLR− NS

XPWB,MXBLR,s = 0 ∀s ∈ SSol (A.118)

Set known heat engine boiler output solid concentrations

xKnHEP,MXBLR,s · NT

HEP,MXBLR− NS

HEP,MXBLR,s = 0 ∀s ∈ SSol (A.119)

Outlet wastewaterUpper bound on output wastewater concentrations

NSMXWW,OUTV,s − xMax

MXWW,OUTV,s · NTMXWW,OUTV

≤ 0 ∀s ∈ SWW (A.120)

Hydrogen/oxygen production

Pressure-swing absorptionSet recovery fraction of H2 from inlet

NSPSA,SPH2P,H2

− RevH2PSA ·

∑(u,PSA)∈UC

NSu,PSA,H2

= 0 (A.121)

Set inlet mole fraction of H2∑(u,PSA)∈UC

NSu,PSA,H2

− InH2PSA ·

∑(u,PSA)∈UC

NTu,PSA = 0 (A.122)

Air separation unitRecovery fraction of O2

NSASU,OUTV,s − (1 − sf ASU) · NS

AC,ASU,s = 0 ∀s ∈ SUASU (A.123)

5 Chem

P

Q

Q

Q

Q

P

F

C

C

C

C

E

C

C

C

L

T

C

C


rocess hot/cold/power utility requirements

Set electricity needed for process units

ElP −

∑u∈UUtil

Su · ElBaseu = 0 (A.124)

Set cooling water needed for process units

CWP −

∑u∈UUtil

Su · CWBaseu = 0 (A.125)

Set heating fuel needed for process units

FCM −∑

u∈UUtil

Su · FBaseu = 0 (A.126)

Set utilities needed for process units

HUu,ut − Su · UBase

u,ut = 0 ∀ut, u ∈ UUtil (A.127)

rocess costs

eedstock costsLevelized cost of biomass feedstock

ostFs =

MWs · NSINBIO,BDR,s · CF

s

Prod · LHVProd∀s ∈ SBio (A.128)

Levelized cost of coal feedstock

ostFs =

MWs · NSINCOAL,CDR,s · CF

s

Prod · LHVProd∀s ∈ SCoal (A.129)

Levelized cost of natural gas feedstock

ostFs =

∑(INNG,u)∈UC

MWs · NSINNG,u,s · CF

s

Prod · LHVProd∀s ∈ SNG (A.130)

Levelized cost of freshwater feedstock

ostFH2O =

MWH2O · NSINH2O,SPWRI,H2O · CF

H2O

Prod · LHVProd(A.131)

lectricity costsLevelized cost of electricity

ostEl = FElIn · CEl

In − FElOut · CEl

Out


O2 sequestration costsLevelized cost of CO2 sequestration

ostSeq =MWCO2 · NS

CO2SC,OUTCO2,CO2

· CSeq


evelized investment costsTotal overnight cost of process units

OCu = (1 + ICu) · (1 + BOPu) · Co,u · Su

So,u

sf u

(A.134)

Variable capital costs of process units

Cu = LCCR · IDCF · TOCu (A.135)

Levelized cost of process units

ostUu = CCu · (1 + OM)

CAP · Prod · LHVProd(A.136)


Objective function

Levelized cost of fuel production

MIN∑u∈UIn

∑(u,s)∈SU

CostFs + CostEl + CostSeq +

∑u∈UInv

CostUu (A.137)

Simultaneous heat and power integration

Pinch pointsSet pinch points based on inlet temperatures

⎧⎪⎪⎪⎪⎪⎨⎪⎪⎪⎪⎪⎩

Tpi = THP-inu,u′ ∀(u, u′) ∈ HP; Tpi = Tu ∀u ∈ HPtHB;

Tpi = Tut ∀(ut, pi) ∈ HPt − PIUt ; Tpi = TPC−inb,c,t

∀(b, c, t) ∈ HEP; Tpi = Tc

Tpi = TCP−inu,u′ + �T ∀(u, u′) ∈ CP; Tpi = TEC−in

b,c+ �T ∀(b, c) ∈ CPEC ;

Tpi = TSH−inb,t

+ �T ∀(b, t) ∈ CPSH ; Tpi = Tut + �T ∀(ut, pi) ∈ CPt − PIUt ;

Tpi = Tb + �T

⎫⎪⎪⎪⎪⎪⎬⎪⎪⎪⎪⎪⎭

(A.138)

Temperature differencesProcess unit hot stream inlets

�THP-inu,u′,pi = max{0, THP-in

u,u′ − Tpi} (A.139)

Process unit hot stream outlets

�THP-outu,u′,pi = max{0, THP-out

u,u′ − Tpi} (A.140)

Process unit cold stream inlets

�TCP−inu,u′,pi

= max{0, TCP−inu,u′ − (Tpi − �T)} (A.141)

Process unit cold stream outlets

�TCP−outu,u′,pi

= max{0, TCP−outu,u′ − (Tpi − �T)} (A.142)

Heat engine precooler inlets

�TPC−inb,c,t,pi

= max{0, TPC−inb,c,t

− Tpi} (A.143)

Heat engine precooler outlets

�TPC−outb,c,t,pi

= max{0, TPC−outb,c,t

− Tpi} (A.144)

Heat engine economizer inlets

�TEC−inb,c,pi

= max{0, TEC−inb,c

− (Tpi − �T)} (A.145)

Heat engine economizer outlets

�TEC−outb,c,pi

= max{0, TEC−outb,c

− (Tpi − �T)} (A.146)

Heat engine superheater inlets

�TSH−inb,t,pi

= max{0, TSH−inb,t

− (Tpi − �T)} (A.147)

Heat engine superheater outlets

�TSH−outb,t,pi

= max{0, TSH−outb,t

− (Tpi − �T)} (A.148)

Heat engine logical existenceBound on heat engine flow rate

FUpb,c,t

· yEnb,c,t ≥ FEn

b,c,t∀(b, c, t) ∈ HEP (A.149)

Bound on total amount of heat engines∑

(b,c,t)∈HEP

yEnb,c,t ≤ EnMax (A.150)

Chem

H

(

Q

L

Q

z

z

�

�


eat balancesHeat engine electricity balance∑

b,c,t)∈HEP

(wTurb,c,t − wPum

b,c,t) · FEnb,c,t = FEl (A.151)

Upper heat balance for pinch points

Hpi =

∑(u,u′)∈HP

∑s

Nsu,u′,s · CpP

u,u′,s · (�THP-inu,u′,pi − �THP-out

u,u′,pi )

+∑

(b,c,t)∈HEP

FEnb,c,t · CpHE−P · (�TPC−in

b,c,t,pi− �TPC−out

b,c,t,pi)

+∑

(ut,pi)∈HPt−PIUt

∑(u,ut)∈HPt

Q HUu,ut +

∑(u,pi)∈HPt−PIHB

Qu

+∑

b

∑(c,pi)∈HPt−PIC

∑t

FEnb,c,t · dHC

c (A.152)

ower heat balance for pinch points

Cpi =

∑(u,u′)∈CP

∑s

Nsu,u′,s · CpP

u,u′,s · (�TCP−outu,u′,pi

− �TCP−inu,u′,pi

)

+∑

(b,c,t)∈HEP

FEnb,c,t · CpHE−E · (�TEC−out

b,c,pi− �TEC−in

b,c,pi)

+∑

(b,c,t)∈HEP

FEnb,c,t · CpHE−S · (�TSH−out

b,t,pi− �TSH−in

b,t,pi)

+∑

(ut,pi)∈CPt−PIUt

∑(u,ut)∈CPt

Q HUu,ut +

∑(b,pi)∈CPt−PIB

∑c

∑t

FEnb,c,t · dHB

b

(A.153)

Pinch point heating deficit

pi = Q Cpi − Q H

pi (A.154)

Negativity of pinch deficits

pi ≤ 0 (A.155)

Total heating deficit

− Qc = 0 (A.156)

Total heat balance

=∑

(u,u′)∈HP

∑s

Nsu,u′,s · CpP

u,u′,s · (THP-inu,u′ − THP-out

u,u′ )

+∑

(b,c,t)∈HEP

FEnb,c,t · CpHE−P · (TPC−in

b,c,t− TPC−out

b,c,t)

+∑

(u,ut)∈HPt

Q HUu,ut +

∑u∈HPtHB

Qu +∑

(b,c,t)∈HEP

FEnb,c,t · dHC

c

−∑

(u,u′)∈CP

∑s

Nsu,u′,s · CpP

u,u′,s · (TCP−outu,u′ − TCP−in

u,u′ )

−∑

(b,c,t)∈HEP

FEnb,c,t · CpHE−E · (TEC−out

b,c− TEC−in

b,c)

−∑

(b,c,t)∈HEP

FEnb,c,t · CpHE−S · (TSH−out

b,t− TSH−in

b,t)

−∑

(u,ut)∈CPt

Q HUu,ut −

∑(b,c,t)∈HEP

FEnb,c,t · dHB

b (A.157)


Appendix B. Supplementary data

Supplementary data associated with this article can befound, in the online version, at http://dx.doi.org/10.1016/j.compchemeng.2012.06.032.

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Proceso Fischer Tropsch

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btu kreutz

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