## Cost and Performance Baseline for
Fossil Energy Plants

Volume 4: Coal

- to

- Liquids via Fischer

Tropsch Synthesis

October 15, 2014

DOE/NETL

- 2011/1477

OFFICE OF FOSSIL ENE RGY



## National Energy Technology LaboratoryCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis



## Disclaimer

This report was prepared as an account of work sponsored by an agency of the United States Government. Neither the United States Government nor any agency thereof, nor any of their employees, makes any warranty, express or implied, or assumes any legal liability or responsibility for the accuracy, completeness, or use fulness of any information, apparatus, product, or process disclosed, or represents that its use would not infringe privately owned rights. Reference therein to any specific commercial product, process, or service by trade name, trademark, manufacturer, or otherwise does not necessarily constitute or imply its endorsement, recommendation, or favoring by the United States Government or any agency thereof. The views and opinions of authors expressed therein do not necessarily state or reflect those of the Uni ted States Government or any agency thereof.

### NETL Contact :
National Energy Technology Laboratory (NETL)

Morgan Summers , General Engineer

Office of Program Performance and Benefits

Prepared By:

Energy Sector Planning and Analysis (ESPA)

Vasant Shah, Norma J. Kuehn, Marc J. Turner

Sheldon J. Kramer

Consultant

This report was prepared by Energy Sector Planning and Analysis (ESPA) for the United States Department of Energy (DOE), National Energy Technology Laboratory (NETL). This work was completed under DOE NETL Contract Number DE

- FE0004001. This work was performed

under ESPA Task 341.02.07.

The authors wish to ac knowledge the excellent guidance, contributions, and cooperation of the NETL staff, particularly:

Thomas J. Tarka

, NETL, Team Lead, OSEAP

Kristin Gerdes

, NETL, Office Director, OPPB

John Wimer

, NETL, Office Director, SCC



### DOE Contract Number DE

- FE0004001This page intentionally left blank.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

i 

## Table of Contents

Executive Summary

................................ ................................ ................................ ......................... 1 Plant Design & Constraints

................................ ................................ ................................ .............. 1

Pe rformance Results

................................ ................................ ................................ ........................ 3

Economic Results

................................ ................................ ................................ ............................. 3 Cost Sensitivity to Financing Structures and Capital Charge Factors

................................ ............. 5

1 Introduction

................................ ................................ ................................ .......................... 7

1.1 Site Description

................................ ................................ ................................ .................... 7

1.2 Design Coal

................................ ................................ ................................ .......................... 8

1.3 Environmental Requirements

................................ ................................ ............................... 8

1.3.1 Carbon Dioxide

................................ ................................ ................................ .................... 9

1.3.2 Mercury

................................ ................................ ................................ .............................. 10

1.4 Raw Water Usage

................................ ................................ ................................ .............. 10

1.5 Balance of Plant

................................ ................................ ................................ ................. 12

1.6 Econ omic Analysis

................................ ................................ ................................ ............ 13

1.6.1 Capital Costs

................................ ................................ ................................ ...................... 14

1.6.2 Production Costs and Expenses

................................ ................................ ......................... 15

1.6.3 Cost of Production of Liquid Fuels

................................ ................................ .................... 16

2 Plant Configuration

................................ ................................ ................................ ............ 18

3 System Description

................................ ................................ ................................ ............ 20

3.1 Synthesis Gas Production

................................ ................................ ................................ .. 21

3.1.1 Air Separation Unit

................................ ................................ ................................ ............ 27

3. 1.2 Coal Drying

................................ ................................ ................................ ........................ 28

3.1.3 Gasifier

................................ ................................ ................................ ............................... 29

3.1.4 Particulate Removal

................................ ................................ ................................ ........... 29

3. 1.5 Gas Scrubbing

................................ ................................ ................................ .................... 30

3.1.6 Water Gas Shift through Acid Gas Removal

................................ ................................ ..... 30

3.1.7 CO 2 Compression and Dehydration

................................ ................................ ................... 33

3.1.8 Claus Unit

................................ ................................ ................................ .......................... 33

3.1.9 Zinc O xide

................................ ................................ ................................ ......................... 33

3.2 Fischer

- Tropsch Synthesis

................................ ................................ ................................ . 34

3.3 Product Recovery

................................ ................................ ................................ ............... 44

3.4 Recycle Gas Loop

................................ ................................ ................................ .............. 45

3.5 Prod

uct Upgrading

................................ ................................ ................................ ............. 45

3.5.1 Naphtha Hydrotreaters

................................ ................................ ................................ ....... 45

3.5.2 Middle Distillate Hydrotreaters

................................ ................................ ......................... 45

3.5.3 Wax Hydrocracking

................................ ................................ ................................ ........... 46

3.6 Power Production

................................ ................................ ................................ ............... 47

3.6.1 Combustion Turbine Generators

................................ ................................ ........................ 47

3.6.2 Steam Turbine Power Generation

................................ ................................ ...................... 49

3.6.3 Flare Stack

................................ ................................ ................................ ......................... 50

4 Performa nce Summary and Mass Balances

................................ ................................ ....... 51

4.1.1 Carbon Balance

................................ ................................ ................................ .................. 53

4.1.2 Sulfur Balance

................................ ................................ ................................ .................... 54

4.1.3 Water Balance

................................ ................................ ................................ .................... 54

5 Cost Estimation Results

................................ ................................ ................................ ..... 56Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

ii 5.1 Capital and Operating Costs

................................ ................................ .............................. 56

5.2 Cost of Production

................................ ................................ ................................ ............. 69

5.3 Cost of Carbon Capture

................................ ................................ ................................ ..... 73

6 Conclusions and Recommendations

................................ ................................ .................. 74

6.1 Conclusions

................................ ................................ ................................ ........................ 74

6.2 Recommendations for Further Study

................................ ................................ ................. 74

7 References

................................ ................................ ................................ .......................... 75Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

iii 

## List of Exhibits

Exhibit ES

- 1 Plant performance summary

................................ ................................ ..................... 3

Exhibit ES

- 2 Capital, operating cost, and COP summary

................................ .............................. 4

Exhibit ES

- 3 TOC ($1,000) with error bars

................................ ................................ .................... 5

Exhibit ES

- 4 Comparison of ECO bbl and EPD bbl

................................ ................................ ...... 6

Exhibit 1

- 1 Site ambient conditions

................................ ................................ ................................ 7

Exhibit 1

- 2 Site characteristics

................................ ................................ ................................ ....... 7

Exhibit 1

- 3 Design coal analysis

– Illinois No. 6

................................ ................................ ........... 8

Exhibit 1

- 4 BACT environmental design basis

................................ ................................ .............. 9

Exhibit 1

- 5 Standards of performance for electric utility steam

- generating units built, reconstructed, or modified after February 28, 2005

................................ ................................ 9

Exhibit 1

- 6 Balance of plant design requirements

................................ ................................ ........ 12

Exhibit 1

- 7 Process and cooling water properties

................................ ................................ ......... 13

Exhibit 1

- 8 Features of an AACE Class 4 cost estimate

................................ .............................. 14

Exhibit 1

- 9 Financial structure for fuels projects

................................ ................................ ......... 16

Exhibit 1

- 10 Financial assumptions for PSFM

................................ ................................ ............. 17

Exhibit 3

- 1 Coal

- to

- liquids process flow diagram (red items only in sequestration case)

........... 20

Exhibit 3

- 2 Block flow diagram of the gasification, syngas shift, and acid gas removal area (sequestration only systems shown in red)

................................ ................................ ........... 22

Exhibit 3

- 3 Stream table for the gasification, syngas shift, and acid gas removal area

– CO 2 sequestration case

................................ ................................ ................................ .................. 23

Exhibit 3

- 4 Stream table for the gasification, syngas shift, and acid gas removal area

– CO 2 vent

case

................................ ................................ ................................ ................................ ........ 25

Exhibit 3

- 5 Typical ASU process schematic

................................ ................................ ............... 28

Exhibit 3

- 6 F

- T reactor syngas operating parameters for both cases

................................ ............ 30

Exhibit 3

- 7 Rectisol process flow diagram example

................................ ................................ .... 32

Exhibit 3

- 8 Simplified representation of the F

- T synthesis section

................................ .............. 35

Exhibit 3

- 9 Block flow diagram for the F

- T synthesis and product upgrading area

.................... 37

Exhibit 3

- 10 Stream table for the F

- T synthesis and product upgrading area

– CO 2 sequestration

................................ ................................ ................................ ................................ ............... 38

Exhibit 3

- 11 Stream table for the F

- T synthesis and product upgrading area

– CO 2 vent

........... 41

Exhibit 3

- 12 Diesel properties

................................ ................................ ................................ ...... 46

Exhibit 3

- 13 Wax hydrocracking model product yield distributions

................................ ............ 46

Exhibit 3

- 14 Block flow diagram for the fuel gas preparation and power generation area

.......... 48

Exhibit 3

- 15 Fuel gas LHV comparison between turbine and F

- T models

................................ .. 49

Exhibit 4

- 1 Plant performance summary

– CO 2 sequestration case

................................ ............. 52

Exhibit 4

- 2 Plant performance summary

– CO 2 vent case

................................ ........................... 53

Exhibit 4

- 3 Elemental carbon mass balance

– CO 2 sequestration case

................................ ........ 54

Exhibit 4

- 4 Elem ental carbon mass balance

– CO 2 vent case

................................ ...................... 54

Exhibit 4

- 5 Water balance

– CO 2 sequestration case

................................ ................................ .... 55

Exhibit 4

- 6 Water balance

– CO 2 vent case

................................ ................................ .................. 55

Exhibit 5

- 1 Total plant cost summary with CO 2 sequestration

................................ ..................... 57

Exhibit 5

- 2 Total operating and maintenance costs with CO 2 sequestration

................................ 62

Exhibit 5

- 3 Total plant cost summary

– CO 2 vent

................................ ................................ ........ 63

Exhibit 5

- 4 Total operating and maintenance costs

– CO 2 vent

................................ ................... 68

Exhibit 5

- 5 Cost estimation results summary (CO 2 sequestration)

................................ ............. 69Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

iv Exhibit 5

- 6 COP sensitivity to market price of electricity

................................ ............................ 70

Exhibit 5

- 7 Comparison of TOC per F

- T bbl, ECO bbl, and EPD bbl

................................ ......... 71

Exhibit 5

- 8 FT

- Liquids COP sensi tivity to CCF

................................ ................................ .......... 72

Exhibit 5

- 9 Cost summary of CTL facility with and without CCS

................................ .............. 73Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

v 

## Acronyms and Abbreviations

AEO Annual Energy Outlook

AGR Acid gas removal

ASTM formerly American Society for

Testing and Materials

ASU Air separation unit

ATR Autothermal reformer

API American Petroleum Institute

BACT Best available control technology

b bl barrel

BEC Bare erected cost

BFD Block flow diagram

BFW Boiler feed water bpd barrel per day

Btu British thermal unit

Btu/hr British thermal unit per hour

Btu/kWh British thermal unit per kilowatt

hour

Btu/lb British thermal unit per pound

CBTL Coal and biomass to liquids

CBM Coal bed methane

CCF Capital charge factor

CCS Carbon capture and sequestration

CF Capacity factor

CFM Cubic feet per minute

CFR Code of Federal Regulations

CH 4 Methane

cm Centimeter

CO Carbon monoxide

CO 2 Carbon dioxide

COP Cost of production of liquid fuels

COR Contracting Officer's Representative

COS Carbonyl sulfide

CT Combustion turbine

CTG Combustion t urbine

- g enerator

CTL Coal

- to

- l iquids

CWR Circulating water return

CWS Circulating water system

CWT Cold water temperature

DCA Direct contact aftercooler

DOE Department of Energy

ECO Equivalent crude oil

EIA Energy Information Administration

EOR Enhanced oil recovery

EPA Environmental Protection Agency

EPC Engineer/ p rocure/ c onstruct EPCM Engineering/ p rocurement/

c onstruction m anagement

E PD Equivalent petroleum diesel

ESPA Energy Sector Planning and

Analysis

F

- T , FT Fischer

- Tropsch

ft Foot, Feet

F TL Fischer

- Tropsch liquid gal Gallon

GE General Electric

GHG Greenhouse gas

G J Giga

joules

G J/hr Giga

joules per hour

gpm Gallons per minute

GT Gas turbine
h, hr Hour

H 2 Hydrogen

H 2 O Water

H 2 S Hyd rogen sulfide

HAP Hazardous air pollutant

HCl Hydrochloric acid
Hg Mercury

HDPE High density polyethylene

HHV Higher heating value hp Horsepower

HP High pressure

HRSG Heat recovery steam generator

HVAC Heating, ventilating, and air

conditioning

IGCC Integrated gasification combined

cycle

IOU Investor owned utility

IP Intermediate pressure

IRR Internal rate or return

IRROE Internal rate of return on equity

ISBL Inside battery limits

ISO International Standards

Organization kg/hr Kilogram per hour

kJ K ilojoules

kJ/hr Kilojoules per hour kJ/kg Kilojoules per kilogram

km Kilometer

kV Kilovolt

kW Kilowatt

kWe Kilowatts electricCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

vi kWh Kilowatt

- hour

kWt Kilowatts thermal

lb Pound

lb/10 6 Btu Pounds per million British thermal

units (also lb/MMBtu)

lb/hr Pounds per hour

lb/ft 2 Pounds per square foot

lb/MMBtu Pounds per million British thermal

units (also lb/10 6 Btu)

lb/MWh Pounds per megawatt hour

LHV Lower heating value

LNB Low NO

x burner

LP Low pressure lpm Liters per minute

m Meters

mg Milligram

m/min Meters pe r minute

m 3 /min Cubic meter per minute

MMBtu Million British thermal units (also

shown as 10 6 Btu)

MMBtu/hr Million British thermal units (also

shown as 10 6 Btu) per hour

MP Medium Pressure

MPa Megapascals

MWe Megawatts electric

MWh Megawatt

- hour

MWt Megawatts thermal

N/A Not applicable

NETL National Energy Technology

Laboratory

N GCC Natural gas combined cycle

NH 3 Ammonia

NOx N itrogen oxides

NSPS New Source Performance Standards

OPP B Office of Program Performance and

Benefits

O&M Operation and maintenance

PM Particulate matter

PM 10 Particulate matter measuring 10 μm

or less

POTW Publicly o wned t reatment w orks

ppm Parts per million

ppmv Parts per million volume ppmvd Parts per million volume, dry PSA Pressure s wing a dsorption

PSFM Power systems financial model psia Pounds per square inch absolute psid Pounds per square inch differential psig Pounds per square inch gage

PT Product transport

RDS Research and Development

Solutions, LLC

RMA Raw material acquisition

RMT Raw material transport

ROE Return on equity

S# Stoichiometric number

SGC Synthesis gas cooler

SGS Sour gas shift

SO 2 Sulfur dioxide

SOx Oxides of sulfur

STG Steam turbine generator

SWS Sour water stripper

TASC Total as

- spent cost

TGTU Tail gas treating unit

TOC Total overnight cost

Tonne Metric t on (1

, 000 kg)

TPC Total plant cost

TPD Tons per day

TPH Tons per hour

TSS Total s uspended s olids

T&S Transport and storage

ULSD Ultra

- low sulfur diesel

U.S. United States

USC Ultra

- supercritical

VGO Vacuum gas oil

vol% Volume percent

WGS Water gas shift

wt% Weight percent

WTI West Texas Intermediate

WTW Well

- to

- Wheels

ZnO Zinc oxide

ZnS Zinc s ulfide
$/bbl Dollars per barrel μS/cm micro Siemens per cm, a measure of

conductivity

 C Celsius

 F FahrenheitCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

1 

## Executive Summary

Economic and national security concerns related to liquid fuels have revived national interest in alternative liquid fuel sources. Coal

- to

- Fischer

- Tropsch fuels production is a major technology option for many states and for the Departme nt of Energy (DOE)

. This report examines the technical and economic feasibility of a commercial coal

- to

- liquids (CTL) facility in the Illinois

coal basin that nominally produces 50,000 barrels per day (bbl/day) F

- T liquids using an iron

based catalyst wit h an expected final product split of approximately 70:30 diesel/naphtha. The facility employs dry

- feed Shell

- type gasification technology and Fischer

- Tropsch (F

- T /FT ) technology to produce commercial

- grade diesel and naphtha liquids from medium

- sulfur

bituminous coal

.
A minimal amount of excess electrical power is also produced and assumed to be sold into the power market to provide an operating credit in the cost estimation calculations

. The scope of this study includes:

- Development of a conceptual p lant design

- Modeling and analysis of the process

- Capital cost estimate

- Economic viability assessment

W here incentives (e.g., loan guarantees) are available , t his study concludes that a CTL facility is economically competitive with crude oil prices as low a s $ 107 per barrel 1

– in the scenario where

CO 2 is not sequestered

– or $ 116 per barrel in the case where CO 2 is sequestered for greenhouse gas ( GHG ) impact mitigation 1

.
As removal of CO 2 from the shifted gas is inher e nt to F

- T operations, a large por tion of the cost of capture is already incorporated into the des ign of a typical F

- T facility. The low incremental cost of sequestration

– about $ 9 per barrel of Equivalent Crude Oil , or $0. 27 per gallon of petroleum diesel equivalent fuel

– demonstrates the relatively low costs of additional capture requirement and compression.

If no incentives are available ( using the commercial fuels financing str ucture ) , the CTL facility is economically viable when the crude oil prices are $ 125 per barrel

– in the scenario where CO 2 is not sequestered

– or $ 1 34 per barrel in the case w he re CO 2 is sequestered for GHG impact mitigation 1

.



### Plant Design & Constraints

The initial plant design used for this study was based on the April 2007 National Energy Technology Laboratory ( NETL ) report

, “Baseline Technic al and Economic Assessment of a

Commercial Scale Fischer

- Tropsch Liquids Facility

. ” [ 1 ] Th e work in this study built upon and incorporated several design changes and improvements that were identified during that analysis

1 The crude oil values listed above are based on Energy Information Administration ( EIA ) data of historic spot prices from June 2009 through November 2013 for the various fuel types. These were used to develop correlations between the various fuel prices an d the West Texas Intermediate (WTI) crude oil price (Crude oil: Ultra

- low sulfur diesel is 1.25 and Naphtha:Diesel is 0.70) . The correlations were used to estimate the Equivalent Crude Oil (ECO)

p rice. The ECO price is the minimum market price for crude o il at which the first

- year COPs will be met.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

2 and through other work

, such as the January 2009 NETL report

, “Affordable, Low

- Carbon

Diesel Fuel fr om Domestic Coal and Biomass

. ” [ 2 ] Design changes from the April 2007 study

include:

1. The gasifier system was changed from the ConocoPhillips E

- GasTM system to a dry

- feed, single

- s tage, entrained

- flow system similar to a Shell Global Solutions commercial design.

2. The operations of the F

- T liquids upgrading units were altered to achieve a product split emphasizing diesel fuel production over naphtha (approximately 70:30). Specificall y, the wax hydrocracking unit was adjusted to provide the desired diesel yield.

3. The process configuration was changed to increase the amount of CO 2 captured for storage by adding a new water

- gas shift (WGS) reactor prior to the combined cycle power generat ion step to convert CO to CO 2 , and by employing a dual

- stage Rectisol ® unit upstream of the F

- T Synthesis loo p for CO 2 separation.

4. A second scenario was added in order to ascertain the incremental cost of CO 2 sequestration. In this scenario, the separated CO 2 is vented to the atmosphere instead of compressing, transporting, and storing it. The CO 2 is still separated to prevent the build

- up of inert CO 2 in the recycle loop and otherwise reduce equipment sizes, and consequently capital costs.

5. The system de sign incorporated low

- cost efforts to reduce water usage.

The CTL plant design that was evaluated utilizes an iron

- based catalyst and a low

- temperature F

T reactor. The concept includes a total of eight gasifiers that produce syngas

, which is cleaned

and combined before being ducted to the F

- T reactor.

Vapor products from the second

- stage reactors (those leaving the F

- T synthesis section) are sent to a cyclone to prevent liquid carryover

. They are then cooled and sent to a three

- phase separator. The ho t condensate is sent to the product upgrading processes. The remaining vapors are processed for CO 2 removal, oxygenate removal, dehydration, and additional hydrocarbon recovery. The majority of the resulting tail

- gas vapor is recycled to the F

- T reactor, with the remainder being processed in the hydrogen recovery unit.

The remaining pre

- fuel gas from the hydrogen recovery unit is mixed with the light gases produced in the hydrocarbon upgrading sections and the off

- gas from the wax stream. In the

case of t he CO 2 sequestration case, this fuel gas is shifted to maximize CO 2 capture potential. In both cases, a portion of the fuel gas is separated out for combustion in direct

- fired heaters, which allows for more steam from the heat recovery to be directed to e lectrical power production. An additional portion of fuel gas is utilized for coal drying and the remainder is sent to the

combustion turbines.

The plant was designed with the goal to operate independent of the bulk electric power system (i.e. the power grid) during steady

- state operations

– producing enough power for plant needs

– while minimizing the production of export power (export power was desired to be less than 10 percent of total electrical power production). The total amount of excess electric al power output varies based on a combination of the internal heat requirements (consumption of fuel gas) and facility heat sources (available for steam production) as well as on whether CO 2 captured at the facility is to be sequestered or merely vented to the atmosphere due to the large electrical load associated with compression.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

3 The electrical power was produced via two sources and in the following ways : 1) the steam from

the F

- T reactor and from the combustion turbine heat recovery unit were combined and utilized in steam turbines, and 2) the aforementioned fuel gas by

- product from the F

- T process was

combusted in off

- the

- shelf GE turbines, which were sized specifically to the available fuel flow in order to achieve the goal of minimizing the export po wer.

The sequestration case achieves the goal of producing less than 10 percent excess electricity (4.6

MW
e ); however, as the vent case was designed to have the same coal flow rate as the sequestration case in order to maintain a liquids output of 50,000 bbl/day , it produces 97.1 MW

e and consequently does not meet the excess electricity production limit.



### Performance Results

Total plant performance data are summarized in Exhibit ES

- 1

. Both facilities were designed to nominally produce 50,000 bpd of liquid fuels, but the CO 2 vent case is slightly more thermally efficient owing to the reduced parasitic load (the result of not having to compress the CO 2 )

.

Exhibit ES

- 1 Plant performance summary

CO 2 Sequestration
Case CO 2 Vent Case

As

- received Coal Feed Rate, tonne/day (tpd) 19,056 (21,006) 19,056 (21,006)

Syngas feed rate, lb/hr 2,498,396 2,498,396

Fuel gas recycle rate, lb/hr 743,921 743,921

Diesel Production, bbl/day 35,230 (70.5%) 35,230 (70.5%)

Naphtha Production, bbl/day 14,762 (29.5%) 14,762 (29.5%)

Gross Plant Power, MW
e 427 473

Net Plant Power for Export, MW
e 4.7 97.1

Carbon Dioxide Captured , tonne/day (tpd) 23,954 (26,405) 23,196 (25,569)

Carbon Dioxide Sequestered, tonne/day (tpd) 23,954 (26,405) n/a

The CTL plant is capable of producing 35,230 bbl/day of commercial

- grade diesel liquid (70.5 percent) and 14,762 bbl/day of F

- T naphtha liquids (29.5 percent). The naphtha would be shipped to a refinery for further upgrading into commercial

- grade end products or for use as a feedstock for the chemicals industry.



### Economic Results

The capital and operating costs, as well as the breakeven selling prices required for the project to be economically viable (meeting a 20 percent Internal Rate of Return on Equity [ IRROE ]), are summarized in Exhibit ES

- 2

. As shown, a commercial

- scale CTL facility will require approximately $ 6.4 and $ 6.5 billion in total overnight cost for the CO 2 vent case and CO 2 sequestration cases, respectively , ranging b etween $ 127,996 and $ 130,662 per barrel of F

- T liquids daily production capacity ($/bbl/day). For comparison purposes, the T OC values are shown in Exhibit ES

- 3 in thousands of dollars ($1,000) with error bars demonstrating the impact

of the +30%/

- 15% accuracy (see S ection 1.6.1 ) in the available cost estimates for this study and consistent with the screening study level of design engineering

.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

4 The primary figure

- of

- merit in this report is the cost of production (COP) (a.k.a. required selling price) expressed in $/bbl of product. The COP values were calculated using the Power Systems Financial Model (PSFM) and estimated to be the value that yields a net present value of zero at a given internal rate of return on equity (IRROE) for 30 years of operation based on the assumed

financial structure and escalations. The COP was assumed to escalate at three percent per year for the thirty

- year economic life of the plant. All costs are expressed in June 2011 dollars. Two

financial structures were assumed for calcula ting the figures

- of

- merit. These structures are based on typical values for fuel projects with and without loan guarantees or government subsidies.

The results are listed in Exhibit ES

- 2

. Two additional figures

- of

- merit included in this study are the equivalent crude oil (ECO) price and the equivalent petroleum diesel (EPD) price. These are the theoretical crude oil and petroleum based diesel price s , based on histor ic data that would produce the calculated COP

values

. Both of the equivalent prices take into account the differences in energy content between

the F

- T and petroleum

- derived fuels (explained further in S ection 1.6.3 ). The se results are also

listed in Exhibit ES

- 2

.
For cases with l oan guarantee incentives, the facility will be economically competitive with traditional (or petroleum) means of fuel production when crude oil prices are between $ 10 7 /bbl and $ 11 6 /bbl for the CO 2 vent and sequestration cases, respectively. This equates to a petroleum

- diesel equivalent price of $ 3.18 /gallon and $ 3.44 /gallon (prices do not include federal or state taxes) for the CO 2 vent and sequestration cases, respectively. The fuel produced in the sequestration case is more expensive due to the compression, transpor t, and storage costs, as evidenced by the positive value for the incremental cost of CCS.

Exhibit ES

- 2 Capital

, operating cost , and COP summary

CO 2 Sequestration CO 2 Vent Incr. CCS
Cost

T otal Overnight Cost , $1000 6,532,052 6,398,760 133,292

Total Overnight Cost , ($/bbl/day) 130,662 127,996 2,666

Total As Spent Cost * , $1000 7,494,877 [7,715,660] 7,341,938

[7,558,216] 152,939

[157,444]
Total As Spent Cost * , ($/bbl/day) 149,922 [ 154,388 ] 146,862 [ 151,189 ] 3,060 [ 3,150 ]

Total annual fixed O&M , $1000 200,8 5 4 197,405 3,448

Total variable O&M (90% CF), $1000 130,076 127,370 2,706
Total annual fuel cost (90% CF), $1000 473,377 473,377 0

Total annual power credit (90% CF), $1000

- 2,149

- 44,844 42,695

COP F

- T Diesel, $/bbl FTD 133.2 [154.2] 123.1 [143.7] 10.1 [10.5]

COP F

- T Naphtha, $/bbl FTN 92.8 [107.4] 85.8 [100.1] 7.0 [7.3]

Equivalent Crude Oil Price*, $/bbl ECO 115.6 [133.8] 106.9 [124.7] 8.7 [9.1]

Equivalent Petroleum Diesel Price*, $/bbl EPD 144.5 [167.2] 133.6 [155.8] 10.9 [11.4]

* Values are shown for two financial structures. The first (lower) value is based on the loan guarantee

finance structure. The second (higher /bracketed

) value is based on the commercial fuels finance

structure.
All costs in June 2011 dollars.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

5 Exhibit ES

- 3 TOC ($1,000) with error bars

Source: NETL



### Cost Sensitivity to Financing Structures and Capital Charge Factors

Two financial structures were examined in this study : t he first structure , considered to be representative of a commercial fuels project

, is assumed to have no loan guarantees

, while the

second structure assumed moderate government incentives

– in the form of government loan

guarantees

. Both scenarios face a high degree of mar ket risk as the facilities involve a large amount of capital that must be financed with revenues from a volatile commodity fuel s market ;

co

- production of electricity via an integrated gasification combined cycle ( IGCC ) plant allows for diversification of p roducts which dampens the impact of the volatile fuels market

. The COP s for both structures are listed in Exhibit ES

- 2

. The costs include a credit of $ 60 /MWh for excess power generated

. T he sensitivity of economic results to the power credit and different financing

structures is included in Section 5.2

.
To achieve e conomic viability, a commercial

- scale CTL facility equipped for CO 2 sequestration will require crude oil prices to be above $133.8 /bbl ( in the case with no incentives ) and $115.6/bbl ( in the moderate incentives case ) to obtain parity with petroleum derived diesel.

Exhibit ES

- 4 summarizes the results of the study, presenting the equivalent crude oil prices and equivalent petroleum diesel prices for each scenario. The sequestration case with no loan

guar a ntees has the highest ECO price and EPD price ( $ 133.8 /bbl and $ 167.2 /bbl, respectively)

. The case that includes incentives and does not sequester CO 2 has the lowest ECO price and EPDCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

6 price ($ 106.9 /bbl and $ 133.6 /bbl, r espectively) due primarily to the lower CCF, but also in part to the increased power production , reduced auxiliary loads, and reduced capital cost resulting

from the exclusion of CO 2 compression systems.

Exhibit ES

- 4 Comparis on of ECO bbl and EPD bbl

Source: NETLCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

7 

## 1 Introduction

The o bjective of this report is to assess the design and financial performance of a commercial

scale coal

- to

- Fischer

- Tropsch liquids plant using a dry

- feed Shell Global Solutions

- type
commercial gasifier with optimum carbon capture.



### 1.1 Site Description

All plants in this study are assumed to be located at a generic plant site in the Midwestern United

States , with ambient conditions and site characteristics as presented in Exhibit 1

- 1 and

Exhibit 1

- 2

.

Exhibit 1

- 1 Site a mbient c onditions

Elevation, ft 0

Barometric Pressure, psia 14.7

Design Ambient Temperature, Dry Bulb,  F 59

Design Ambient Temperature, Wet Bulb,  F 51.5

Design Ambient Relative Humidity, % 60

Exhibit 1

- 2 Site c haracteristics

Location Greenfield, Midwestern United States

Topography Level

Size, acres 300

Transportation Rail , Road, Pipeline

Ash Disposal Off Site

Water Municipal (50%) / Groundwater (50%)

Access Landlocked

, with access by railway and highway

CO 2 Storage Delivered to plant boundary at 15.2 MPa (2,215 psia)

The following design parameters are considered site

- specific and are not quantified for this study. Allowances for normal conditions and construction are included in the cost estimates.

 Flood plain considerations

 Existing soil/site conditions

 Water discharges and reuse

 Rainfall/snowfall criteria

 Seismic design

 Buildings/enclosures

 Fire protection

 Local code height requirements

 Noise regulations , and their i mpact on site and surrounding areaCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

8 

### 1.2 Design Coal

The design coal assumed for this study is presented in Exhibit 1

- 3

.

Exhibit 1

- 3 Design coal analysis

– Illinois No. 6

Rank Bituminous

Seam Illinois No. 6 (Herrin)

Source Old Ben mine [ 3 ]

Proximate Analysis (weight %) a

As Received Dry

Moisture 11.12 0.00

Ash 9.70 10.91

Volatile Matter 34.99 39.37

Fixed Carbon 44.19 49.72

Total 100 .00 100 .00

Sulfur 2.51 2.82

HHV, kJ/kg 27,113 30,506

HHV, Btu/lb 11,666 13,126

LHV, kJ/kg 26,151 29,544

LHV, Btu/lb 11,252 12,712

Ultimate Analysis (weight %)

As Received Dry

Moisture 11.12 0.00

Carbon 63.75 71.72

Hydrogen 4.50 5.06

Nitrogen 1.25 1.41

Chlorine 0.29 0.33

Sulfur 2.51 2.82

Ash 9.70 10.91

Oxygen b 6.88 7.75

Total 100 .00 100 .00

Notes:

a The proximate analysis assumes sulfur as volatile matter

b By difference



### 1.3 Environmental Requirements

The environmental control equipment used in the conceptual design conform s to b est a vailable

c ontrol t echnology (BACT) guidelines

, which are summarized in Exhibit 1

- 4

.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

9 Exhibit 1

- 4 BACT environmental design basis

Pollutant Environmental Design Basis

Control Technology Limit

Sulfur Oxides (SO 2 ) Rectisol ® / Selexol/Econamine
Plus/Sulfinol

- M + Claus Plant 99 + % or ≤ 0.0 50 lb/10 6 Btu

Nitrogen Oxides (NOx) LNB and N 2 Dilution primarily
with humidification as needed 15 ppmvd (@ 15% O 2 )

Particulate Matter (PM) Cyclone/Barrier Filter/Wet
Scrubber/AGR Absorber 0.0 06 lb/10 6 Btu

Mercury (Hg) Activated Carbon Bed 95% removal

While the coal

- to

- liquids ( CTL ) process is considered a petrochemical complex, the regulations on emissions considered in this study were for fossil

- fuel

- fired power plants. The current regulation governing new fossil

- fuel

- fired power plants is the New Source Performance Standards (NSPS) published in February 2006 and shown in Exhibit 1

- 5 . These sta ndards supersede the previous performance standards established in 1978

. [ 4 ] The new standards apply to units with the capacity to generate greater than 73 MW of power by burning fossil fuels, as well as cogeneration units that sell more than 25 MW of pow er and more than one

- third of the potential output capacity to any utility power distribution system. The rule also applies to combined cycle, including integrated gasification combined cycle ( IGCC ) plants, and combined heat and power combustion turbines that burn 75 percent or more synthetic

- coal gas.

Exhibit 1

- 5 Standards of performance for electric utility steam

- generating units built,

reconstructed, or modified after February 28, 2005

Pollutant New Units Reconstructed Units Modified Units

Emission

Limit %

Reduction Emission

Limit (lb/10 6 Btu) %

Reduction Emission

Limit (lb/10 6 Btu) %

Reduction

PM 0.015 lb/10 6 Btu 99.9 0.015 99.9 0.015 99.8

SO 2 1.4 lb/MWh 95 0.15 95 0.15 90

NOx 1.0 lb/MWh N/A 0.11 N/A 0.15 N/A

The BACT emission limits assumed for this study exceed the emission requirements of the 2006 NSPS. It is possible that state and local requirements could supersede NSPS or BACT and impose even more stringent requirements.

1.3.1 Carbon Dioxide

For both the sequestration and vent cases, the basis is an optimum removal based on carbon input

in the coal

, and excluding carbon in the Fischer

- Tropsch ( F

- T ) product and the carbon that exits the gasifier with the slag. CO 2 is captured from the syngas l eaving the gasifiers using the

Rectisol ® process, from the vapor stream exiting the F

- T reactor using a proprietary amine CO 2 - removal system , and in the sequestration case only CO 2 is captured from the f uel gasCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

10 downstream of the water gas shift ( WGS ) unit using a proprietary amine system. The design includes compression of the recovered CO 2 to 2,215 psia for pipeline transport at the plant boundary. The CO 2 is transported 100 km ( 62 miles) via pipeline to a geologic sequestration field for injection into a saline formation. The CO 2 is transported and injected as a supercritical

fluid in order to avoid two

- phase flow and achieve maximum efficiency. CO 2 transport and

storage (T & S ) performance specifications were independently estimated by the Nationa l Energy Technology Laboratory ( NETL )

. The T&S estimation methodology is explained in more detail

in the Quality Guidelines for Energy System Studies (QGESS) document “ Carbon Dioxide Transport and Storage Costs in NETL Studies

. ” [ 5 ]

1.3.2 Mercury

The mercury content in the Illinois No. 6 coal is reported as a mean value of 0.09 ppm (dry basis)

, with a standard deviation of 0.06

, based on 34 samples of coal shipped by Illinois mines. Mercury capture from the syngas via activated carbon is assumed to be 95 perc ent

, based on experience at the Eastman Chemical plant. If additional capture is necessary, dual carbon beds can be used to achieve greater than 99 percent mercury removal.



### 1.4 Raw Water Usage

Raw water makeup is assumed to be provided 50 percent by publicly owned treatment works (POTW) and 50 percent from groundwater. The amount of raw water usage for an IGCC case has been estimated in a previous study done for NETL

. [ 5 ] In that study, raw water usage was defined as the water metered from a raw water source and used in the plant processes for any and all purposes, such as cooling tower makeup, condenser makeup, ash handling makeup, syngas

humidification

, and quench system makeup. In that study, as in the current study, the plant is equipped with an evaporative cooling tower, and all process blowdown streams are treated and recycled to the cooling tower. This is the “lowest hanging fruit” on the water minimization

s ystem.

Since the F

- T plant differs from an IGCC plant, the technique for estimating water usage is

different and based on the following :

1. The power generation processes w ere modeled in the Aspen process simulation program

, which calculates the required cooling water load and stack water loss.

2. The Aspen model for the F

- T synthesis area and the product upgrading area w as used to predict the required F

- T cooling water load.

3. The total of the cooling tower blowdown and vaporization losses was predicted based on the estimated cooling water circulation rate.

Other water minimization options are possible but were not factored into the present study. One option to reduce the makeup water demand would be to add a reverse osmosis unit and treat the cooling tower sy stem blowdown streams and recycle them either to the cooling tower or steam systems. However, most of the water loss is evaporative loss from the cooling towers. More air

coole d heat exchangers could be added to reduce the load on the cooling tower and its corresponding evaporative losses.

Another water minimization option includes recovering water from the coal dryer. The Shell

based gasification process uses a dry

- feed system

, which is sensitive to the coal moisture content.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

11 Coal moisture consists of two parts : surface moisture and inherent moisture. For coal to flow smoothly through the lock hoppers, the surface moisture must be removed. The Illinois No. 6 coal used in this study contains 11.12 percent total moisture on an as

- received basis. It was assumed that the coal must be dried to 5 percent moisture to allow for smooth flow through the dry

- feed system. The coal is simultaneously crushed and dried in the coal mill then delivered to a surge hopper with an approximate 2

- hour capacity. The dr ying medium is provided by combining the off

- gas from the Claus plant and a slipstream of clean syngas and passing them through an incinerator. The incinerator flue gas, with an oxygen content of 6 volume percent, is then used to dry the coal in the mill. The design does not include the recovery of water from the coal drying system.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

12 

### 1.5 Balance of Plant

Assumed balance of plant requirements are listed in Exhi bit 1

- 6

.

Exhi bit 1

- 6 Balance of plant design requirements

Cooling system Recirculating, Evaporative Cooling Tower Cooling tower makeup water composition is available

. [ 3 ]

Fuel and Other storage

Coal 30 days

Slag 30 days

Sulfur 30 days

Plant Distribution Voltage

Motors below 1 hp 110/220 volt

Motors 250 hp and below 480 volt

Motors above 250 hp 4,160 volt

Motors above 5,000 hp 13,800 volt

Steam and Gas Turbine generators 24,000 volt

Grid Interconnection voltage 345 kV

Water and Waste

Makeup Water Process water is available at a flow rate of 1,500 gpm

. [ 3 ]

Feed water Process water treatment is included and will produce boiler feed quality water.

Process Wastewater Process wastewater and storm water that contacts equipment
surfaces is collected and treated for recycle as slurry or quench makeup. Selected blowdown is discharged through a permitted discharge permit.

Sanitary Waste Disposal Design include s a packaged domestic sewage treatment plant with effluent discharged to the indust rial wastewater treatment system. Sludge is hauled off site.

Water Discharge Most of the wastewater is recycled for plant needs. Blowdown is treated for chloride and metals, and discharged.

Solid Waste S lag is assumed to be a solid waste that is classified as non

hazardous. An offsite waste disposal site is assumed to have the capacity to accept waste generated throughout the life of the facility. Solid waste sent to disposal is at an assumed nominal fee per ton, even if the waste is hauled back t o the mine. Solid waste generated that can be recycled or reused is assumed to have a zero cost to the technology

.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

13 Process water and cooling water come from two different treatment facilities. The composition and physical properties of each are shown in Exhibit 1

- 7

.

Exhibit 1

- 7 Process and cooling water properties

Property Process Water Cooling Water

Total Dissolved Solids (TDS) 200 μS/cm 1250 μS/cm

Total Suspended Solids (TSS) N/A N/A

Hardness 100 mg/l as CaCO 3 75 mg/l as CaCO 3 Alkalinity 100 ppm 350 ppm

Sulfate 4 ppm 50 ppm

Chloride 10 ppm 200 ppm

Silica 30 mg/l 30 mg/l

Aluminum N/A N/A

Iron 0.25 mg/l 0.25 mg/l

Calcium 70 mg/l 25 mg/l

Magnesium 25 mg/l 45 mg/l

Phosphate 0.4 mg/l 6.0 mg/l (ortho)

Ammonia <1 mg/l 19 mg/l

Chlorine <0.1 mg/l <0.1 mg/l

pH 8.0 8.0



### 1.6 Economic Analysis

Estimates of capital cost, operating costs

, and financial performance were updated using the best available information and expressed in June 20 11 dollars.

Capital cost and production cost estimates developed for the CTL plant were based on adjusted

vendor

- furnished, actual cost data, or best possible projections (when necessary). The financial performance of the facility was estimated using various cost s , product prices

, and financial

parameters to generate a c ost of production ( CO P) value.

Detailed information pertaining to topics such as contracting strategy, EPC contractor services, estimation of capital cost contingencies, owner’s costs, cost estimate scope, economic assumptions, finance structures, cost of electricity, etc. are available in the April 2011 revision of the QGESS document “Cost Estimation Methodology for NETL Assessment of Power Plant Performance” [ 6 ] Select portions are repeated here for completeness.

Despite the application of process contingencies and other be st efforts to realistically project the cost of emerging technologies, it should be noted that such cost projections are often shown to be optimistic when compared to actual costs. Accordingly, the "next commercial offering" cost estimates in this report, while intended for plants that lie somewhere between first

- of

- a

- kind (FOAK) and nth

- of

- a

- kind (NOAK) in the maturity spectrum, could be closer to NOAK plant costs given the inherent "technology optimism" associated with projecting costs for emerging

tech nologies. Continuing research and development could result in technologies that are more advanced than the emerging technologies assessed by this report, leading to costs that are lower

than those estimated herein.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

14 Detailed information regarding FOAK and NOAK can be found in the August 2013 revision of the QGESS document “Technology Learning Curve (FOAK to NOAK)” [ 7 ]

1.6.1 Capital Costs

The cost estimation methodology is explained in more detail in Section 2.7 of “Cost and Performance Baseline for Fossil Energy Power Plants, Volume 1: Bituminous Coal and Natural Gas to Electricity

. ” [ 8 ]

Bare erected costs (BEC) comprises the cost of process equipment, on

- site facilities and infrastructure that support the plant (e.g., shops, offices, labs, road), and the direct and indirect labor required for its construction and/or installation. The cost of EPC services and contingencies are not included in BEC. BEC is an overnight cost expressed in base

- year dollars.

Engineering, procurement, and construction costs (EP CC) comprise the BEC plus the cost of services provided by the engineering, procurement and construction (EPC) contractor. EPC services include: detailed design, contractor permitting (i.e., those permits that individual contractors must obtain to perfo rm their scopes of work, as opposed to project permitting, which is not included here), and project/construction management costs. EPCC is an overnight cost expressed in base

- year dollars.

Total plant costs (TPC) comprise the EPCC plus project and process contingencies. TPC is an overnight cost expressed in base

- year dollars.

Total overnight capital costs (TOC) comprise the TPC plus all other overnight costs, including owner’s costs. TOC is an “overnight” cost, expressed in base

- year dollars and as suc h does not include escalation during construction or interest during construction. The TOC for each plant was calculated by adding owner’s costs to the total plant cost (TPC).

Total as

- spent capital costs (TASC) are the sum of all capital expenditures a s they are incurred during the capital expenditure period including their escalation. TASC also includes interest during construction. Accordingly, TASC is expressed in mixed, current

- year dollars over the capital expenditure period.

The capital costs ha ve an estimated accuracy of + 30 /

- 15 percent, consistent with the screening study level of design engineering applied to the two cases in the study. The accuracy reflects the

r ecommended practice 18R

- 97 guidelines of the AACE (formerly Association for the Advancement of Cost Engineering International ) for estimates intended for the purpose of a “Feasibility Study” (AACE Class 4). [ 9 ] Exhibit 1

- 8 describes the character istics of an AACE

Class 4 cost estimate. The value of the study lies not in the absolute accuracy of the individual cases, but in the fact that both cases were evaluated under the same set of technical and economic assumptions. The consistency of approac h allows meaningful comparisons.

Exhibit 1

- 8 Features of an AACE Class 4 cost estimate

Project
Definition Typical Engineering Completed Expected Accuracy

1 to 15%  plant capacity, block schematics, indicated layout, process

flow diagrams for main process

 systems, and preliminary engineered process and utility

 equipment lists - 15% to

- 30% on the low side, and +20% to +50% on the high sideCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

15 The costs are grouped according to a process/system oriented c ode of accounts. This type of

code

- of

- account structure has the advantage of grouping all reasonably allocable components of a system or process so they are included in the specific system account. (This would not be the case had a facility, area, or com modity account structure been chosen instead).

Process contingency was added to cost account items that were deemed to pose significant risk due to lack of operating experience. Process contingency is intended to compensate for uncertainty in cost estim ates caused by performance uncertainties associated with the development status of a technology. Process contingencies are applied to each plant section based on its current technology status using guidelines in the AACE International Recommended

Practice 16R

- 90. [ 10 ]

Project contingency was added to the EPCC accounts to cover the cost of any additional equipment or other costs that would result from a more detailed design. Project contingencies

are added to the summation of EPC and process contingency c osts.

Both contingencies represent costs that are expected to occur. Each BEC account was evaluated against the level of estimate detail and field experience to determine project contingency.

TOC values are expressed in J une 20 11 dollars. The estimate represents current commercial offerings for the gasification and F

- T technologies

, or best possible projections for very near

- term,
yet non

- commercial

, offerings. The estimates represent a complete fuels plant facility. The

bounda ry limit is defined as the total plant facility within the “fence line,” including the coal receiving facilities , power generation facility, and water supply system. Costs were grouped according to a process/system

- oriented code of accounts

.

1.6.2 Production Costs and Expenses

The production or operating and maintenance (O&M) costs described in this report section pertain to charges associated with operating and maintaining the entire plant over its expected life.

O&M costs are determined on a n annual basis fo r the first year of operation . Quantities for major consumables such as fuel were taken from the heat and mass balance developed for this application. Other consumables evaluated on the basis of the quantity required the use of reference data. Operating labor costs were determined on the basis of the number of operators.

Maintenance costs were evaluated on the basis of requirements for each major plant section.

The O&M costs and expenses associated with the plant include:

 Operating labor

 Maintenance

– material and labor

 Administrative and support labor

 Consumables

 Fuel cost

 Taxes and insurance

These costs and expenses are estimated on a reference basis and escalated to June 20 11 dollars.

The costs assume normal operation and do not include the initial startup costs. The operating labor, maintenance material and labor, and other labor

- related costs w ere combined and then divided into two components: f ixed O&M costs , which are independent of liquids production,Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

16 and variable O&M costs , which are proporti onal to liquids production. The variable O&M cost estimate allocation is based on the plant capacity factor.

The other operating costs, consumables and fuel, are determined on a daily 100

- percent operating capacity basis and adjusted to an annual plant op eration basis. The inputs for each category of operating costs and expenses are identified in the succeeding subsections, along with more specific discussion of the evaluation processes.

The petroleum equivalent diesel price is calculated by taking the co st of production for diesel in $/bbl and multiplying this value by the ratio of the lower heating values of FT diesel and petroleum diesel.

Petroleum Equivalent Diesel Price = ( Petroleum Diesel LHV

FT Dies el LHV )

∗ FY COP FT Diesel

The equivalent crude oil price is then calculated by multiplying the petroleum equivalent diesel price by a factor of 0.80. The factor of 0.80 was calculated from data of historic spot prices provided by the EIA f

rom June 2009 through November 2013 for various fuel types.

Equivalent Crude Oil = 0

. 80 ∗ Petroleum Equivalent Diesel Price

This data was used to develop correlations between the various fuel prices and the WTI crude oil price (Crude oil: Ultra

- low sulfur diesel is 1.25 and Naphtha: Diesel is 0.70). The ECO price is the theoretical crude oil price, based on historic data

, which would produce the calculated COP.

1.6.3 Cost of Production of Liquid Fuels

The figure

- of

- merit in this report is the CO P expressed in $/ bbl . The CO P s were calculated using the Power Systems Financial Model (PSFM) which is a discounted cash flow model

. [ 11 ]

The COP values is estimated to be the required revenue that yields a net pres ent value of zero at a given internal rate of return on equity (IRR OE ) for 30 years of operation based on the assumed

financial structure and escalations. COP is assumed to escalate at three percent per year for the thirty

- year economic life of the plant. All costs are expressed in June 20 11 dollars

. The capital and operating costs for CO 2 T&S were modeled based on the specifications in the NETL QGESS: [ 5 ] In the sequestration case, the costs were combined with the plant capital and operating costs to produce an overall COP.

The re are two project financial structure s: one is representative of commercial fuel projects and the second is representative of a project with loan guarantees . The economic assumptions used

to estimate the COP s are shown in Exhibit 1

- 9 and Exhibit 1

- 10

. [ 12

, 13 ]

Exhibit 1

- 9 Financial structure for fuels projects

Scenario Commercial Fuels Loan Guar a ntees

Percent Debt 50 % 60 %

Percent Equity 50% 40%

Debt Interest Rate 8 .00 % 4.56%

Internal Rate of Return on Equity (IRROE) 20% 20%

After Tax Weighted Cost of Capital 12.48% 9.70%

Capital Charge Factor (CCF) 0.218 0.170Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

17 Exhibit 1

- 10 Financial a ssumptions for PSFM

Parameter Value

TAXES

Income Tax Rate 38% (Effective 34% Federal, 6% State) Capital Depreciation 20 years, 150% declining balance

Investment Tax Credit None Tax Holiday None

FINANCING TERMS Repayment Term of Debt 30 years Grace Period on Debt Repayment None

Debt Reserve Fund None

TREATMENT OF CAPITAL COSTS Capital Cost Escalation During Construction (nominal annual rate) 3.6 % a

Distribution of Total Overnight Capital over the Capital Expenditure Period (before escalation) 5

- Year Period: 10%, 30%, 25%, 20%, 15%

Working Capital zero for all parameters

% of Total Overnight Capital that is Depreciated 100% (this assumption introduces a very

small error even if a substantial amount of TOC is actually non

- depreciable)

INFLATION COP, O&M, Fuel Escalation (nominal annual rate) 3.0% b COP , O&M, ECO

, Fuel

a A nominal average annual rate of 3.6 percent is assumed for escalation of capital costs during

construction. This rate is equivalent to the nominal average annual escalation rate for process plant construction costs between 1947 and 2008, according to the Chemical Engineering Plant Cost Index.

b An average annual inflation rate of 3.0 percent is assumed. This rate is equivalent to the average annual escalation rate between 1947 and 2008 for the U.S. Department of Labor's Producer Price
Index for Finished G oods, the so

- called "headline" index of the various Producer Price Indices. (The Producer Price Index for the Electric Power Generation Industry may be more applicable, but that data does not provide a long

- term historical perspective since it only dates back to December 2003.)Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

18 

## 2 Plant Configuratio n

The CTL process configuration chose n for this study is designed to

(1) Maximize the conversion of feedstocks to liquid transportation fuels, achieved through recycling unreacted synthesis gas back to the reactor;

(2) Maximize diesel fuel as the primary liquid fuel product;

(3) Produce enough electrical power for the facility to be self

- sufficient , but minimize exported electrical powe r to a level no greater than 10 percent of the internal power requirements for the CO 2 se questration case ;

(4) Minimize the amount of CO 2 and other greenhouse gases released to the atmosphere during production of the fuel, in so far as plant performance or final product cost is not significantly impacted.

The facility was designed to produce 5 0,000 bbl/day of F

- T liquids using an iron

- based catalyst

with an expected final product split of approximately 70:30 diesel/naphtha. This is in contrast to the April 2007 study wherein F

- T naphtha represented a more significant portion of the final produ ct, based on a desire to focus on gasoline production (via the upgrading of naphtha).

The production of additional diesel fuel in the current configuration requires that extra hydrogen (compared to the previous design) be supplied to the diesel hydroproc essing units in the product upgrading step

. In addition, the wax hydrocracker model was redesigned to ach ieve h igher diesel fuel yields. The diesel produced in the facility is ready for delivery to end

- use customers, after the addition of the appropriate additives (additives are added at the distribution site) . The naphtha produced has been stabilized and can either be shipped to a refinery for further upgrading to a commercial grade gasoline blendstock or sold as a chemical feedstock.

The plant achieves the goal of electrical power independence, in part, by recovering heat from

the Fischer

- Tropsch synthesis reactors to produce steam, which is then used in a steam turbine along with steam produced by syngas cooling and in the heat recovery steam generatio n unit. In

the CO 2 sequestration case, the steam derived power represents 26 percent (29 percent in the CO 2 vent case) of the electrical power generated in the CTL facility with the remaining 74 percent (71 percent in the CO 2 vent case) being produced by the combustion of light gases recovered from the reactor vapor overhead and the various product recovery and upgrading sections of the CTL facility in the combustion turbines.

The final product slate of the facility is nominally :

 35,230 bbl/day of F

- T dies el fuel

 14,76 2 bbl/day of stabilized (heteroatom, olefinic, and aromatic content is reduced) F

- T naphtha

 The sequestration case produces 427.2 MWe of electrical power, 422.5 MWe of which is used within the facility and 4.7 MWe of which is sold into the grid ( 1.1 percent of total

power production at a credit of $60/MWh . The vent case produces 472.8 MWe of electrical power, 375.7 MWe of which is used within the facility and 97.1 MWe of which is sold into the grid (20.5 percent o f total power production ) at a credit of $60/MWh

.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

19  The sequestration case captures 26, 405 tons/day of CO 2 (2, 200,423 lb/hr) and compress es it to approximately 2,215 psia for pipeli ne transport from the facility to a geologic

formation 100 km (62 miles) away

. The vent case vents 25,569 tons/day of CO 2 ( 2, 200,424 lb/hr) from the capture systems.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

20 

## 3 System Des cription

The CTL plant design utilized in this report represents the culmination of “lessons learned” over (a) the course of a number of years studying CTL facilities, (b) shifting environmental and performance concerns, and (c) numerous sensitivity analyses to ascertain optimal process performance.

A simplified process flow diagram of the major unit processes is provided in Exhibit 3

- 1 in order to provide an overview of the CTL plant configuration. The CO 2 vent case differs only in that it

does not include CO 2 compression , water gas shift of the fuel gas, or CO 2 capt ure from the fuel

gas.

Exhibit 3

- 1 Coal

- to

- l iquids process flow diagram (red items only in sequestration case)

Source: NETL

Synthesis gas (syngas) is produced in a dry

- feed S hell type gasifier

. In order for the syngas to be

usable in an F

- T synthesis reactor, it must be cleaned and treated. The major equipment used to clean the syngas consists of a water

- wash scrubber for particulate removal (which also has the added benefit of removing some of the mercury), an activated carbon bed for mercury removal, a WGS reactor 2 , a dual

- stage Rectisol ® acid gas removal (AGR) system using a refrigerated

methanol solution for sulfur and CO 2 removal 3

. A Claus plant is used to produce elemental

sulfur from the sulfur

- containing gases captured in the AGR system.

2 Only a portion of the syngas is passed through the WGS reactor. The amount of syngas shifted is controlled to

obtain an overall H 2 :CO molar ratio between 1:1 and 1.1:1, as required for the F

- T process.

3 CO 2 is removed to reduce the volume of syngas, and therefore equipment sizes and capital expenditures. H 2 S and COS are removed in order to protect the F

- T catalyst.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

21 The clean syngas is then mixed with tail gas recycle from the hydrogen recovery unit and fed to

the low

- temperature, slurry

- phase, iron

- based, F

- T synthesis reactor, which converts the syngas to saturated hydrocarbons. The majority of the hydrocarbon products (those with a carbon chain length greater than four) are upgraded in the hydrocracking and hydrotreating units (Product Upgrading in the f

low diagram) to produce F

- T diesel (70 percent by volume) and F

- T naphtha (30 percent by volume). The tail gas from the F

- T unit, which includes unconverted CO and H 2 , short chain hydrocarbons, and CO 2 is processed in an amine unit for CO 2 removal. This prevents
the buildup of CO 2 in the recycle loop and reduces CO 2 emissions from the facility (in the scenario where captured CO 2 is sequestered). A portion of the tail gas is then directly recycled to

the F

- T reactor for maximum conversion of carbon to fue ls, while the remainder is sent to: (a) a combustion turbine and subsequent heat recovery steam generator (HRSG), (b) direct fired heaters for coal drying, and (c) direct fired heaters for process heating needs.

The sequestration and vent cases differ in that:

1) In the sequestration case the CO 2 streams from the syngas capture system (Rectisol ® ), F

T tail gas capture system (amine), and F

- T fuel gas capture system (amine) are combined and compressed, transported, and sequestered in a geologic formation. 2) In the vent case, there is no compression, transportation, or sequestration. The CO 2 is still captured in the syngas capture system and F

- T tail gas capture system but not in the

F

- T fuel gas capture system. The CO 2 produced from the two capture systems is vented into the atmosphere.

The steam produced by the recovery of heat in the F

- T synthesis reactor, along with the steam produced in the HRSG, is utilized in a steam turbine to produce power.

The combined output of the combustion and steam turbines is wi thin 10 percent of total plant electricity requirement for the CO 2 capture case

. Individual process components are described in greater detail in the following section s

.



### 3.1 Synthesis Gas Production

Exhibit 3

- 2 provide s a simplified block flow diagram of the gasification and syngas cleanup sections. The primary differences between the CO 2 vent and sequestration cases are that: (a)

CO 2 is captured in the CO 2 vent case, but it is not compressed, but instead vented, (b) the CO 2 vent case does not have water gas shift of the fuel gas or an amine CO 2 capture system for the fuel gas , and (c) the absence of the se fuel gas treatment systems and CO 2 compression resu lts in

a lower auxiliary power requirement , as well as lower cooling water requirement for the facility, therefore, the amount of power exported increases and the amount of water withdrawal

decreases

. These differences are shown in the performance summari es provided in Exhibit 4

- 1

and Exhibit 4

- 2

.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

22 Exhibit 3

- 2 Block flow diagram of the gasification, syngas shift, and acid gas removal area (sequestration only systems shown in red)

Source: NETLCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

23 Exhibit 3

- 3 Stream table for the gasification, syngas shift, and acid gas removal area

– CO 2 sequestration caseCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

24 Exhibit 3

- 3 Stream table for the gasification, syngas shift, and acid gas removal area

– CO 2 sequestration case (continued)Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

25 Exhibit 3

- 4 Stream table for the gasification, syngas shift, and acid gas removal area

– CO 2 vent caseCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

26 Exhibit 3

- 4 Stream table for the gasification, syngas shift, and acid gas removal area

– CO 2 vent case ( c ontinued)Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

27 3.1.1 Air Separation Unit

The facility is equipped with a total of four air separation units ( ASU ) for the eight gasifier trains. Each ASU, composed of two cold boxes, is designed to produce a nominal output of 15,182 tonne/day (16,739 tpd) of 95 mole percent O 2 . This O 2 is used in the gasifier, F

- T plant, and Claus plant. The ASU is designed with two production trains. The air compressors are

power ed by an electric motor. There is no direct integration between the ASU and the rest of the facility. A process schematic of a typical ASU is shown in Exhibit 3

- 6

The air feed to the air separation unit is supplied from stand

- alone air compressors. The filtered air is compressed in centrifugal compressors, with intercooling between each stage. The ai r stream is cooled and then fed to an adsorbent

- based pre

- purifier system.

Air from the stand

- alone compressor is cooled and fed to an adsorbent

- based pre

- purifier system. The adsorbent removes water, CO 2 , and C 4 + saturated hydrocarbons in the air. Aft er passing through the adsorption beds, the air is filtered with a dust filter to remove any adsorbent fines that may be present. Downstream of the dust filter a small stream of air is withdrawn to supply the instrument air requirements of the ASU.

Regene ration of the adsorbent in the pre

- purifiers is accomplished by passing a hot nitrogen stream through the off

- stream bed(s) in a direction countercurrent to the normal airflow. The nitrogen is heated against extraction steam (1.7 MPa [250 psia]) in a shel l and tube heat exchanger. The regeneration nitrogen drives off the adsorbed contaminants. Following regeneration, the heated bed is cooled to near normal operating temperature by passing a cool nitrogen stream through the adsorbent beds. The bed is re

- pressurized with air and placed on

stream so that the current on

- stream bed(s) can be regenerated.

The air from the pre

- purifier is then split into three streams. About 70 percent of the air is fed directly to the cold box. About 25 to 30 percent of the air is compressed in an air booster compressor. This boosted air is then cooled in an aftercooler against cooling water before it is

fed to the cold box. About 5 percent of the air is fed to a turbine

- driven, single

- stage, centrifugal booster compressor. This stream is cooled in a shell and tube aftercooler against cooling water

before it is fed to the cold box.

All three air feeds are cooled in the cold box to cryogenic temperatures against returning product oxygen and nitrogen streams in plate

- and

- fin heat exchangers. The large air stream is fed directly to the first distillation column to begin the separation process. The second air stream is liquefied against boiling liquid oxygen before it is fed to the distillation columns. The third,

small air s tream is fed to the cryogenic expander to produce refrigeration to sustain the cryogenic separation process. The work produced from the expansion is used to power the turbine booster

compressor.

Inside the cold box the air is separated into oxygen and nit rogen products. The oxygen product is withdrawn from the distillation columns as a liquid and is pressurized in a cryogenic pump. The pressurized liquid oxygen is then vaporized against the high

- pressure air feed before being warmed to ambient temperatur e. The gaseous oxygen exits the cold box and is split into two streams. Essentially all of the gaseous oxygen is fed to the centrifugal compressor with intercooling between each stage of compression. The compressed oxygen is then fed to the gasification unit. The second oxygen stream is fed to the Claus plant, the autothermal reformers

in the F

- T area, and to the autothermal reformer in the fuel gas cleanup area.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

28 Nitrogen is produced from the cold box at two pressure levels. Low

- pressure nitrogen is sp lit into two streams. The majority of the low

- pressure nitrogen is compressed and fed to the gas turbine as diluent nitrogen. A small portion of the nitrogen is used as the regeneration gas for the

pre

- purifiers and recombined with the diluent nitrogen. A high

- pressure nitrogen stream is also produced from the cold box and is further compressed before it is also supplied to the gas

turbine.

Exhibit 3

- 5 Typical ASU p rocess s chematic

Source: NETL

3.1.2 Coal Drying

The Shell

- based gasification process uses a dry

- feed system that is sensitive to the coal moisture content. Coal moisture consists of two parts: surface moisture and inherent moisture. For coal to flow smoothly through the lock hoppers, the surfac e moisture must be removed. The Illinois No. 6 coal used in this study contains approximately eleven weight

- percent total moisture on an as

received basis. It was assumed that the coal must be dried to approximately 5 percent moisture to allow for smooth flow through the dry feed system.

The coal is simultaneously crushed and dried in the coal mill and then delivered to a surge hopper, which has approximately a two

- hour capacity. The drying medium is provided by

c ombining the off

- gas from the Claus plant and a slipstream of F

- T tail gas and passing themCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

29 through an incinerator. The incinerator flue gas, with an oxygen content of 6 volume percent, is used to dry the coal in the mill.

The coal is drawn from the surge hoppers and fed through a pressurized lock hopper system to a dense phase pneumatic conveyor, which uses CO 2 recovered from the syngas to convey the coal to the gasifiers.

The equipment in the coal drying and grinding system is fabricated of materials app ropriate for the abrasive environment present in the system. The tanks and agitators are rubber lined. The pumps are either rubber

- lined or hardened metal to minimize erosion. Piping is fabricated of high

- density polyethylene (HDPE).

3.1.3 Gasifi er

The gasifi er is a dry

- feed, single

- stage, entrained

- flow gasification system similar to a Shell Global Solutions commercial design. The gasifier system includes high

- temperature gas cooling via syngas recycle, slag handling, and solid waste handling equipment. The gasifier also includes a refractory

- lined water wall that is protected by molten slag that solidifies on the

cooled walls.

The heat recovery from the gasifier is broken into two stages. The first stage 1) recovers high

temperature heat in each gasifier t rain via a gasifier jacket, which cools the syngas by maintaining the reaction temperature at 1,427°C (2,600°F). The second stage 2) reduces the temperature of the product gas from the gasifier to 1,082°C (1,980°F) by mixing in recycled gas,

then to 399 ° C (750°F) with water quench, which provides a portion of the water required for the

WGS reaction.

The gasification system consists of four gasifier trains with two gasifiers in each train, along with one spare gasifier, for a total of nine dry

- feed gasif iers . The spare gasifier is included to maintain overall plant availability of 90 percent. The gasification plant section was modeled as four copies of a single gasification train to enable direct comparison with baseline study case

models of similar siz e, and the results were combined to generate the required quantity of syngas

for the F

- T facility.

These gasifiers, operate at 4.2 MPa (615 psia) and process a total of 19,056 tonne/day (21,006 tpd) of as

- received coal for both the CO 2 sequestration and ve nt cases. Coal reacts with oxygen and steam at a temperature of 1,427 °C (2,600 °F) to produce principally H 2 and CO,

with little CO 2 or CH 4 formed.

3.1.4 Particulate Removal

After passing through the water quench , the syngas passes through a cyclone and a raw gas candle filter system where a majority of the fine particles are removed and returned to the gasifier with the coal fuel. The filter consists of an array of ceramic candle elements in a pressure vessel. Fines produced by the gasification system are re circulated to extinction. The ash that is not carried out with the gas forms slag and runs down the interior walls, exiting the gasifier in liquid form. The slag is solidified in a quench tank for disposal. Lockhoppers reduce the pressure of the solids from 4.2 MPa to atmospheric pressure 1.0 MPa (615 to 14.7 psia).

Following the water quench and particulate removal, the syngas is cooled to 232°C (450°F) before going to the syngas scrubber by vaporizing high

- pressure boiler feed water (BFW) andCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

30 pre

- hea ting intermediate

- pressure BFW. The syngas scrubber removes residual particulate

matter further downstream.

3.1.5 Gas Scrubbing

The “sour” gas leaving the particulate filter system consists mostly of H 2 , CO, CO 2 , water vapor,

N 2 , and smaller quantities of CH 4 , COS, H 2 S, and NH 3

. The quench scrubber washes the syngas in a counter

- current flow in two packed beds. The syngas leaves the scrubber saturated at a temperature of 193 ° C ( 379 oF). The quench scrubber removes essentially all of the entrained particles, p rincipally unconverted carbon, slag, and metals. The bottoms from the scrubber are sent to the slag removal and handling system for processing.

Sour water from the syngas scrubber is stripped of sour gas and treated for recycle or discharge.

The sour w ater stripper removes chlorides, NH 3 , and other impurities from the waste water stream of the scrubber. The sour gas stripper consists of a sour drum that accumulates sour water from the gas scrubber and condensate from synthesis gas coolers. Sour water from the drum flows to the sour water stripper, which consists of a packed column with a steam

- heated reboiler. Sour gas is stripped from the liquid at about 250 °F and sent to the sulfur recovery unit. Remaining water is sent to wastewater treatment.

3.1.6 Wa ter Gas Shift through Acid Gas Removal

The required operating parameters of the F

- T synthesis reactor are satisfied with the WGS and AGR units. These parameters (H 2 :CO, H 2 O:C O , and stoichiometric number [ S# ] ) are described below and reported in Exhibit 3

- 6

. The metrics apply to the total F

- T reactor feed, which includes the clean syngas, steam and recycle from the hydro carbon recovery process.

Exhibit 3

- 6 F

- T reactor syngas operating parameters for both cases

Parameter Value

S# 1 0.98

H 2 :CO 2 1.02

H 2 O:CO 3 1.80

1 Synthesis gas metric at the inlet of the F

- T reactor

2 Ratio of hydrogen to carbon monoxide at the inlet of the F

- T reactor

3 Ratio of moisture to carbon monoxide at the inlet to the WGS units

The H 2 :CO ratio is simply calculated by dividing the moles of hydrogen by the moles of carbon monoxide in the syngas stream. Due to the ability of the iron

- ba sed F

- T catalyst to promote water gas shift in situ enabling CO 2 to be a source of carbon , the F

- T reactor requires a relatively

low H 2 /CO molar ratio of 1:1 to 1.1:1.

The H 2 O:CO ratio was based on common practice and assures an atmosphere conducive to shifting and provides a suitable moisture content at the WGS exit to prevent coking

.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

31 The S# metric is calculated using the following equation:

S # = ( H 2 − CO 2 ) ( CO + CO 2 )

⁄

In systems with out significant water gas shift activity, the S# is targeted to valu es near 2 [ 14 ]

. Since the iron

- based catalyst promotes the WGS reaction, the S# is less than 2 in the study cases

. The process descriptions are provided in the following subsections.

3.1.6.1 Water Gas Shift

In order to achieve the desired H 2 :CO ratio of 1:1 at the reactor inlet , about 15 percent of the syngas is routed through two WGS reactors in series. The WGS reactors convert most of the CO

to H 2 and CO 2 by reacting the CO with water over a fixed bed of catalyst via the wate r gas shift

reaction shown bel ow.

Water Gas S h ift : CO + H 2 O ↔ C O 2 + H 2 Steam is injected prior to the first stage of WGS to adjust the H 2 O/CO molar ratio.

As water gas shift is a net exothermic process, the heat of reaction must be managed between stages. The exiting temp erature of the first stage is 470 ° C (878 oF), which is an increase over the inlet temperature of 219 ° C (426 oF). The temperature is reduced by utilizing a heat exchanger that reduces the temperature to 204 ° C (400 oF) by producing steam.

The temperature again rises to 254 ° C (489 oF) in the second stage of WGS. This heat is

recovered (the temperature is reduced to 230 ° C (446 oF)) by preheating the syngas exiting the gas scrubber (described in Section 3.1.5 ) to 198 °C ( 389 °F) from 193 °C (379 °F) prior to entering the first stage of WGS.

The two stages of WGS result in a 95.6 percent overall conversion of CO to CO 2 . The WGS catalyst also serve s to hydrolyze COS thus eliminating the need for a separate COS hydrolysis

reactor.

The shifted syngas rejoins the bypassed syngas and enters the low temperature cooling system,

where it is cooled to 35 °C (95 °F) prior to the mercury removal beds.

3.1.6.2 Mercu ry Removal

Mercury is removed by packed beds of sulfur

- impregnated carbon similar to what has been used at the Eastman Chemical gasification plant in Kingsport, Tennessee. A bed of sulfur

impregnated carbon, with approximately a 20

- second superficial gas residence time, can achieve

95 percent mercury removal in addition to removal of other volatile heavy metals, such as arsenic. The spent carbon sorbent is discarded as a hazardous waste with a disposal cost of $1,300/ton

.
3.1.6.3 Acid Gas Removal

A feature of this plant configuration is that COS, H 2 S, and CO 2 are removed from the syngas within the same process system, the Rectisol ® unit. The purpose of the Rectisol ® unit, which

utilizes a refrigerated methanol solvent, is to preferentially remove H 2 S and COS a s a product stream, and then to remove the remaining H 2 S and a majority of the CO 2 . This is achieved in theCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

32 dual

- stage Rectisol ® unit. Since the Rectisol ® system removes COS, the COS hydrolysis pre

treatment step is unnecessary and has been eliminated.

The use of methanol prevents formation of ice and hydrates at the low operating temperatures. Cool, dry, particulate

- free syngas enters the first absorber unit at approximately 100 oF and

520 psia. The low operating temperatures are beneficial, because at these temperatures the solvent carrying capacity for both CO 2 and H 2 S becomes very high. (1) These features allow for high separations of both CO 2 and H 2 S.

The low temperature operation can lead to complicated flow schemes , as well as elevated plant

costs. (1) Therefore, Rectisol generally finds application only in relatively difficult gas treating conditions where extremely high purities are required.

Shown below in Exhibit 3

- 7 is an example Rectisol process employed as the Texaco refinery, in Wilmington, CA. (2) The unit was built by Linde A.G. of Munich and is part of a hy drogen production train where vacuum residual oil is gasified to provide hydrogen for the refinery.

Exhibit 3

- 7 Rectisol process flow diagram example

Source: EPA

The feedgas (stream 1) is first cooled against the separation products, and is sent to the main

absorber where H 2 S and CO 2 are removed in the bottom section of the tower. The sulfur

- free gas is washed in the top section of the main absorber to remove remaining CO 2 ( stream 2). The bottom product from the main absorber contains all H 2 S, some CO 2 , and some H 2 . This stream is

flashed to recover H 2 from the loaded methanol solvent. The H 2 is pressurized and recycled to the feedgas inlet. The bottoms product of the flash is sent to the H 2 S enrichment column to

remove additional CO 2 to achieve the required H 2 S concentration. This is accomplished with the stripping gas (stream 6), which is nitrogen obtained from the oxygen plant. The CO 2 rich off gas (stream 3b) is combine d with the off gas from the cold stripper (stream 3a) and is the product

CO 2 stream (stream 3), which is vented to the atmosphere or compressed (see S ection 3.1.7 )

.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

33 The bottom of the H 2 S enrichment column is sent to the warm stripper where the absorbed H 2 S

and CO 2 are stripped and sent to the Claus plant (stream 4).

The stream of CO 2 - loaded methanol, leaving the CO 2 - absorption section of the main absorber (middle s ection) after H 2 flashing, is stripped with nitrogen in the cold stripper, and the overhead (stream 3a) is combined with stream 3b, as discussed prior. The bottoms of the cold stripper, partially stripped, cold methanol, is sent back to the main absorber. The methanol/water mixture condensed from the feed gas is separated in the methanol water separation still to produce an overhead methanol stream, which is recycled, and a bottoms waste water stream (stream 5).

The amine CO 2 capture systems associated wi th the F

- T tail gas and F

- T fuel gas are discussed in

S ection 3.2

.

3.1.7 CO 2 Compression and Dehydration

CO 2 is generated at three pressure levels from the AGR process. In the CO 2 vent case, this CO 2 is vented to the atmosphere, while in the sequestration case the low pressure ( LP ) stream is compressed from 0.15 MPa (22 psia) to 1.1 MPa (160 psia) and then combined with the medium pressure ( MP ) stream. The high pressure ( HP ) stream is combined between compressor stages at 2.1 MPa (300 psia). The combined stream is compressed from 2.1 MPa (300 psia) to a supercritical condition at 15.3 MPa (2,215 psia) using a multiple

- stage, intercooled compressor. During comp ression, the CO 2 stream is dehydrated to a dewpoint of

- 40 oC (

- 40 °F) with triethyleneglycol. The raw CO 2 stream from the Rectisol ® process contains over 9 9 percent CO 2 with the balance being primarily nitrogen and hydrogen , f or modeling purposes it was treated as 100 percent CO 2

. The pure CO 2 is transported via pipeline to a geologic sequestration site 100 km (62 miles)

.

3.1.8 Claus Unit

Acid gas from each Rectisol ® stripper unit is routed to the associated Claus plant. The Claus plant partially oxidizes the H 2 S in the acid gas to elemental sulfur. The rate of sulfur recovery is approximately 43,875 lb/hr in both cases. This rate represents ~100 percent recovery.

Acid gas from the Rectisol ® unit is preheated to 450 °F. A portion of the acid gas, along with al l of the sour water stripper overhead vapor and oxidant, is fed to the Claus furnace. In the furnace,

H 2 S is oxidized to SO 2 using 95 percent oxygen. A furn ace temperature greater than 2,450 °F must be maintained in order to thermally decompose all of the NH 3 present in the sour gas

stream.

Three preheaters and three sulfur converters are used to obtain a per

- pass H 2 S conversion efficiency of approximately 97 .8 percent. In the furnace waste heat boiler, 650 psia steam is generated. This steam is used to satisfy all Claus process preheating and reheating requirements as well as to produce excess steam that is sent to the medium

- pressure steam header. The sul fur condensers produce 50 psig steam that goes to the low

- pressure steam header.

3.1.9 Zinc Oxide

While the Rectisol ® system is capable of reducing the sulfur levels to between 0.1 and 1.0 ppm, which are sufficiently low for an F

- T process using iron

- based catal yst, the long term on

specification operation of the Rectisol ® unit cannot be guaranteed.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

34 Therefore, t he treated, clean syngas is fed to a packed bed of solid zinc oxide (ZnO) to remove any residual heteroatom content (i.e.

, sulfur) that makes it through the syngas treatment processes. The zinc oxide is a final guard bed to protect the Fischer

- Tropsch synthesis catalyst and is capable of reducing the sulfur content of the sweet gas to the F

- T feed to less than 1 ppb.

The cost of the addition of the requi red two ZnO beds to the process is insignificant compared to the total plant cost. If the sulfur polishing step were not there during an upset in the Rectisol ®

unit, the entire amount of F

- T catalyst in the reactors might be poisoned and would have to be replaced. The cost of the shutdown and lost production would far outweigh the cost of the sulfur polishing unit. The addition of this unit is insurance in the same manner that spares are installed for almost all liquid pumps.



### 3.2 Fischer

- Tropsch Synthesis

The cleaned syngas is preheated and mixed with steam and recycled tail gas and fed to the slurry

- bed F

- T hydrocarbon synthesis reactor (which operates at around 475 o F and 300 psia)

. The coal

- derived syngas has a hydrogen

- to

- carbon monoxide (H 2 :CO) ratio o f 0.73, prior to tail gas recycle (1.02 post tail gas recycle), which is converted in the F

- T reactor to straight chain olefinic and aliphatic hydrocarbons with a H:C ratio of approximately 2.1. Use of an iron

- based

F

- T catalyst (as opposed to cobalt

- base d) is preferred when utilizing low H 2 :CO ratio syngas as

the iron

- based catalyst also promotes WGS activity, which allows a portion of the WGS reaction to occur in the reactor. This additional WGS activity within the reactor produces a liquid product with a H:C ratio that is higher than the H 2 :CO ratio of the inlet syngas.

The F

- T reactor consists of two parallel first

- stage slurry bed s that feed a common second

- stage slurry bed . The conceptual design of the F

- T synthesis section is represented in Exhibit 3

- 8

. Unconverted synthesis gas is cooled between the first and second stages to 65.6 o C (150 o F) to remove water and a hot hydrocarbon condensate (which is sent to pr oduct recovery), and then re

heated prior to entering the second stage. Each stage accomplishes approximately 61 percent conversion of carbon monoxide (hydrogen is in excess) leading to a per

- pass conversion of 85 percent of carbon monoxide. There is no direct recycle of unconverted synthesis gas back into

the second

- stage reactor (the implementation of the recycle loop is discussed in a subsequent section).

Reactor temperature is modulated by removing tremendous heat released by the Fischer

- Tropsch

proce ss through generation of 24.8 bar ( 360 psig ) steam through bayonet tubes suspended from

an internal double tubesheet. The reactors operate at 21

- 22 bar (305

- 319 psia) and nominally

249 o C (480 o F) 4

. Synthesis gas is introduced into the reactors through a cylindrical gas distributor and bubbles up through the slurry phase. Fischer

- Tropsch wax and catalyst make up the slurry phase, into which the synthesis gas dissolves, contacts the catalyst, and is oligomerized, producing more Fischer

Tropsch liquids. St reams of liquid wax are withdrawn from the reactors with a slurry valve, which accomplishes the majority of the catalyst

- wax separation.

4 The outlet temperature is set by the target that was yield in the systems model, with this c hoice being at the edge of the operating window referred to as low temperature Fischer

- Tropsch synthesis.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

35 Exhibit 3

- 8 Simplified representation of the F

- T synthesis section

S ource: NETL

Vapor is removed from the overhead of the first

- stage reactor passing through a cyclone to prevent liquid carryover. The vapor is cooled and sent to a three

- phase separator. The hot condensate (light liquid hydrocarbons) is recovered for furt her use. The water is sent to treatment, and the remaining vapors are reheated and sent to the second

- stage reactor.

A cyclone removes entrained catalyst particles from the vapor stream leaving the top of the F

- T reactor before being sent to a condenser f or recovery of condensable hydrocarbon products, which are then processed in hydrotreaters in the product upgrading section of the plant.

The remaining vapor stream

– containing unreacted syngas, inerts, and light hydrocarbons (predominantly methane, ethan e, and propane)

– is processed in an amine unit for CO 2 removal to prevent CO 2 build up in the recycle loop and reduce equipment sizes. In the sequestration case, the captured CO 2 is sent to the gasification section for compression with CO 2 from the

Recti sol ® unit. In the vent case it is merely vented to the atmosphere. A portion of the vapor stream is compressed for recycle to the F

- T reactor while the rest of the gas is sent through a hydrogen recovery process to separate out hydrogen that is used for hydrotreating the liquids. The remaining gas is used as fuel gas.

The wax products are sent to hydrocyclones to accomplish the remaining catalyst

- wax separation. Subsequently, the resulting wax stream is degassed with the resulting vapor cooled

to remove additional condensate that is sent to product recovery and the residual vapor combined with the tail gas for recycling. The liquid wax is sent to wax c l arifying and fed to the hydrocarbon recovery section.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

36 The catalyst that is separated from the bulk of the wax streams is recycled back to the Fischer

Tropsch reactors. To maintain performance, a portion of the catalyst inventory of the reactors is removed for disposal and replaced. Catalyst replacement rates are influenced by loss of catalys t functional lifetime (activity and selectivity) and attrition losses. A composite lifetime (mechanical and functional) of the catalyst in this system analysis is 0.78 years, meaning 0.015 percent of catalyst inventory is replace d per hour.

The catalyst

- f

ree liquid leaving the filter system is reduced in pressure and flashed. The vapor stream is further cooled to 100 °F and flashed. The vapor stream from the second flash is used as fuel for the gas turbine and the liquid from both flash drums is sent to the feed drum of the product fractionators, along with the condensate recovered from the overhead vapor streams.

Exhibit 3

- 9 provides simplified block flow diagram of the F

- T, product recovery, and product upgrading areas. Exhibit 3

- 10 and Exhibit 3

- 11 show the temperature, pressure, flow, and compositions of the process streams in the F

- T area at the design coal feed rate for the CO 2 sequestration and CO 2 vent cases, respectively.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

37 Exhibit 3

- 9 Block flow diagram for the F

- T synthesis and product upgrading area

Source: NETLCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

38 Exhibit 3

- 10 Stream t able for the F

- T synthesis and product upgrading area

– CO 2 s equestrat ionCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

39 Exhibit 3

- 10 Stream t able for the F

- T synthesis and product upgrading area

– CO 2 s equestration (continued)Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

40 Exhibit 3

- 10 Stream t able for the F

- T synthesis and product upgrading area

– CO 2 s equestration (continued)Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

41 Exhibit 3

- 11 Stream t able for the F

- T synthesis and product upgrading area

– CO 2 v entCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

42 Exhibit 3

- 11 Stream t able for the F

- T synthesis and product upgrading area

– CO 2 v ent ( c ontinued)Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

43 Exhibit 3

- 11 Stream t able for the F

- T synthesis and product upgrading area

– CO 2 v ent ( c ontinued)Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

44 

### 3.3 Product Recovery

All of the hydrocarbon streams from the Fischer

- Tropsch synthesis se ction (wax and condensate recovered from the overhead vapor streams) are combined in the feed drum of the product fractionator. The product fractionator separates these hydrocarbons into a low

- pressure light hydrocarbon vapor stream and three liquid strea ms: naphtha, middle distillate, and wax. The liquid cuts of the fractionators are defined by the following boiling range cut points:

Naphtha C 5 saturates to 350°F (177°C)

Middle Distillate 350

- 650°F (177

- 343°C)

Wax Greater than 650°F (343°C)

This simple separation section consists of one fractionation column with a steam side

- stripper. Required support equipment (heat exchangers, flash drums, and furnace utilities) are included in the capital estimate of this section. The feed drum operates as a flash drum at approximately 135°F (57°C) and 90 psia (6.2 Bar). The vapor feed rate to the tower is used for pressure control of the flash drum. The vapor serves as stripping media after being fed to the middle of the column. The liquid is heated to app roximately 343°F (173°C) via a fired heater, which burns residual fuel gas.

The naphtha product is withdrawn from an upper tray and is sent to the naphtha hydrotreating section. Steam stripping is used to provide a stable liquid middle distillate stream fr om the middle of the column, which is cooled (via contact with the column feed stream), and then is sent to the distillate hydrotreating section. The tower bottoms are withdrawn and cooled via contact with the tower feed and subsequent water

- cooled exchang ers prior to being sent to the wax hydrocracking section.

The Hydrogen Recovery system processes a slipstream of the F

- T tail gas to recover H 2 for use in the hydrotreaters and hydrocracker. A pressure swing adsorption (PSA) unit is used to produce high pu rity H 2 and the rejected stream from the process is blended with the remaining F

- T tail gas for use as fuel gas.

The rationale for placement of the hydrogen recovery downstream of the F

- T reactor is three

fold:

1. The presence of H 2 promotes the F

- T reaction:

CO + 2n H 2 = CnH 2 n + n H 2 O.

2. The F

- T reactor produces additional H 2 via the WGS reaction:

CO + H 2 O = CO 2 + H 2

. 3. Less gas is treated per unit of H 2 recovered since the F

- T liquids are removed.

Fuel gas recovered from the Hydrocarbon Recovery, Hydrogen Reco very, Naphtha Hydrotreating, Distillate Hydrotreating, and Wax Hydrocracking units is mixed, flashed, and then compressed to 350 psia at 110 ° F.

Oxygen and steam are then added to this stream prior to a WGS reactor in order to increase the

H 2 and CO 2 fra ctions in the fuel gas. After the WGS reactor in the sequestration case , this stream is treated by proprietary amine AGR process which is used to capture CO 2 , thereby reducing the greenhouse gas ( GHG ) emissions in the turbine exhaust. This captured CO 2 is

mixed with the CO 2 removed from the F

- T synthesis loop and sent to the gasification island forCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

45 dehydration and compression. In the vent case the fuel gas is combusted without shifting or CO 2 capture

. The remaining fuel gas at 335 psia is separated and part of it routed to the various F

- T process furnaces. The remaining F

- T fuel gas is sent to the gasification area for coal drying and power production. A simplified block flow diagram of the power generation area of the facility is provided in S ection 3.6.1

.



### 3.4 Recycle Gas Loop

Vapor products from the second

- stage reactors (thos e leaving the Fischer

- Tropsch synthesis section) are sent to a cyclone to prevent liquid carryover, cooled, and then sent to a three

- phase separator. The hot condensate is sent to the product upgrading processes. The remaining vapors are processed for CO 2 removal

, oxygenate removal, dehydration, and additional hydrocarbon recovery. The majo rity of the resulting tail

- gas vapor is recycle d to the F

- T reactor, with the remainder being processed in the hydrogen recovery unit.

The remaining pre

- fuel gas from the hydrogen recovery unit is mixed with the light gases produced in the hydrocarbon upgr ading sections and the off

- gas from the wax stream . In the

case of the CO 2 sequestration case, this fuel gas is shifted to maximize CO 2 capture potential. In both cases, a portion of the fuel gas is separated out for combustion in direct fired heaters throughout the CTL plant, which allows for more steam from the heat recovery to be directed to electrical power production, and the remain der is sent to the combustion turbines.



### 3.5 Product Upgrading

3.5.1 Naphtha Hydrotreaters

Light liquid hydrocarbon products are processed in the naphtha hydrotreating units to produce stabilized (heteroatom, olefinic, and aromatic content is reduced) F

- T naphtha. T his is then blended with naphtha produced in the hydrocracker. The off gas from this process is combined

wi th portions of the F

- T tail gas.

3.5.2 Middle Distillate Hydrotreaters

The product recovery distillation tower produces a middle distillate that is sent t o a fixed bed catalytic hydrotreater that produces a n acceptable quality diesel fuel. This processing section removes oxygenates and any residual heteroatoms (sulfur and nitrogen), and greatly reduces the

olefinic and aromatic content. The distillate hydr otreater also produces, and the processing section separates off, a waste

- water stream and a residual C4

- vapor stream, which is recovered. The hydrotreated middle distillate from this processing section is blended with the analogous

stream from the wax h ydrocracker to form a product diesel. The product diesel has an estimated pour point of

- 45.6°C (

- 50.1°F), which is well below the target value of any region in the United States [ 15 ]

, and the estimated cetane index of this diesel product is 70.7, which is above the standards set in the World Wide Fuel Charter (minimum requirements between 48 and 55 depending on category [ 16 ]); hence there is no need for additives. These values are provided in

Exhibit 3

- 12

.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

46 Exhibit 3

- 12 Diesel properties

Property Value

Density, lb/bbl 266.9

Net Heat of Combustion, MM Btu /gal (LHV) 0.12

Pour Point, °F

- 50.1

Cetane Index 70.7

3.5.3 Wax Hydrocracking

The models utilized in this study are based on Chevron’s isocracking process that emphasized diesel production compared to naphtha production

. [ 1 0] Based on the reported yields for a two

stage design processing VGO (Vacuum Gas Oil), a set of wax hydrocracking yields was developed. Exhibit 3

- 13 provides the wax hydrocracking model product yield. The wax hydrocracking model produces a Diesel/(C5+ naphtha) product ratio of 6.

Exhibit 3

- 13 Wax hydrocracking model product yield distributions

Componen t Yields

Weight % Volume %

C 1 - C 4 s 4.50

-

C 5 - C 6 s 4.59 5.58

Naphtha 6.67 7.39

Diesel 84.24 87.03

Total 100.00 100.00

C 5 +Naphtha 11.26 12.97

Diesel/(C 5 +Naphtha) 7.48 6.71

The F

- T process design can be adjusted to a low wax yield (as little as 56 weight percent) and

still make the desired 2.3 3 /1 Diesel/(C5+Naphtha) ratio by using the wax hydrocracking model described previously.

This facility was designed for maximum diesel fuel production and the hydrocracker is configured to maximize diesel produ ction by selectively hydrocracking the wax product mainly to the diesel boiling range. The wax fraction is sent to the wax hydrocracker, which breaks down the wax into naphtha and middle distillate, as well as some residual C4

- vapor, which is

recovered. The single, multibed, downflow wax hydrocracker catalytic ally cracks the wax under

a 75

- 103 bar ( 1088

- 1494 psia ) hydrogen atmosphere at a nominal temperature of 371°C ( 700°F )

. There is inter

- bed cooling to remove the heat of reaction from hydrogen

- rich rec ycle gas. A flash drum is used to recover unutilized hydrogen

- rich gas for recycle back to the hydrocracker. A distillation column is used to separate the liquids from the flash into naphtha, middle distillate,

and wax streams. The residual uncracked wax is recycled to extinction. The middle distillate from the wax hydrocracker is sent for blending with the middle distillate range product from theCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

47 distillate hydrotreating section. The naphtha stream from the wax hydrocracker is sent to the naphtha reformin g unit.

To achieve a 70:30 diesel/naphtha product split (volume basis), the wax hydrocracker model was redesigned from that used in the April 2007 study to increase the diesel fuel yield. In the process, a certain amount of hydroisomerization occurs, whi ch has the beneficial effect of lowering the pour point of the diesel fuel.



### 3.6 Power Production

The plant is designed to produce sufficient electric power to satisfy the power requirements for the gasification plant and the F

- T facility, to make the entire fa cility electric power independent. Limiting the amount of excess power produced has t wo benefits: 1 ) the facility is not dependent on external power source s or supplies and 2) it improves the efficiency of the facility.

The prevailing industry philosophy in designing CTL facilities is that excess syngas should only

be combusted in a direct

- fired boiler so as to maximize the amount of syngas converted to the higher value liquid fuel products. It was determined, however, that the use of a combined

- cycle

po wer cycle (i.e.

, adding a combustion turbine to the facility) improved both the environmental and economic performance of the facility, and subsequently it was decided to use such a power cycle. Commercially available combustion turbines of the appropriat e size were selected i n order to minimize excess power production.

3.6.1 Combustion Turbine Generators

The combustion turbine generators that could be used for this application are three General Electric (GE) gas turbines (MS6001B) each producing a nominal 45 MWe. The actual total power production is 112 MWe in the CO 2 sequestration case and 136 MWe in the CO 2 vent case. The difference in power output between the cases is attributed to differences in fuel gas composition . The stream data are available in Exhibit 3

- 3 and the block diagram is provided in

Exhibit 3

- 14Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

48 Exhibit 3

- 14 Block flow diagram for the fuel gas preparation and power generation area

Source: NETLCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

49 Because fuel gas from the F

- T area is imported into the gasification and power production A spen model, some adjustment had to be made to the composition of the imported fuel gas. The A spen gasification and power plant model do not contain any hydrocarbons heavier than methane. Therefore, in order to import the fuel gas into the power plant model, the heavier hydrocarbons were changed to methane on a weight basis. This overestimates the heating value and tends to increase the power production because methane has a higher heating value than the heavier hydrocarbons. However, this change is slight s ince there is only a relatively small amount of these hydrocarbons in the total fuel gas going to the gas turbine. Using the sequestration case as an example, Exhibit 3

- 15 shows there is negligible effect on lower heating value (LHV) by substituting trace heavy components with methane.

Exhibit 3

- 15 Fuel g as LHV comparison between turbine and F

- T m odels

Component LHV

Btu/lb Mass Fraction

in F

- T Fuel Gas LHV F

- T Fuel

Gas , Btu/lb Mass Fraction

in Turbine Fuel

Gas LHV Turbine

Fuel Gas

, Btu/lb

H 2 51,573 0.02301009 1186.69937 0.0230101 1186.69937

CO 4,344 0.0818232 355.43998 0.0818232 355.43998

CO 2 0 0.00211792 0.00000 0.0021179 0.00000

H 2 O 0 0.00296592 0.00000 0.0029659 0.00000

NH 3 7,998 0.00025819 2.06500 0.0002582 2.06500

HCN 9,915 8.51E

- 07 0.00844 8.51E

- 07 0.00844

CH 4 21,509 0.40201125 8646.85998 0.4021163 8649.11842

C 2 H 4 20,275 1.18E

- 06 0.02394 0 0.00000

C 2 H 6 20,424 0.00010435 2.13124 0 0.00000

C 3 H 6 19,675 8.44E

- 09 0.00017 0 0.00000

C 3 H 8 19,920 6.39E

- 08 0.00127 0 0.00000

IC 4 H 8 19,469 1.16E

- 11 0.00000 0 0.00000

10193.22939 10193.33121

3.6.2 Steam Turbine Power Generation

Heat from the various gasification, AGR, and F

- T processes is recovered with multiple heat exchangers matching the quality and quantity of available heat. This heat recovery produces steam at various pressures, part of which is used for process needs and the remainder is used to run the steam turbines to produce additional power of 316 MWe and 337 MWe of power in the sequestration and CO 2 vent cases, respectively.

The HRSG is a horizontal gas flow, drum

- type, multi

- pressure design that is matched to the

ch aracteristics of the gas turbine exhaust gas. The HP drum produces steam at main steam pressure; while the intermediate pressure ( IP ) drum produces steam for export to the cold reheat. The HRSG drum pressures are nominally 1,800 /452 psia for the HP/IP tu rbine sections, respectively.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

50 Natural circulation of steam is accomplished in the HRSG by utilizing differences in densities due to temperature differences of the steam. The natural circulation HRSG provides the most

cost

- effective and reliable design.

3.6.3 Fl are Stack

Each gasifier has a self

- supporting, refractory

- lined, carbon steel flare stack to combust and dispose of product gas during startup, shutdown, and upset conditions. The flare stack is provided with multiple pilot burners, fueled by natural gas or propane, with pilot monitoring

instrumentation.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

51 

## 4 Performance Summary and Mass Balances

T he plant power output, including auxiliary load, for the facilit ies is summarized in Exhibit 4

- 1

and Exhibit 4

- 2

.
The total F

- T liquids production of 49,992 bpd consists of 35,23 0 bpd (70.5 percent ) diesel and 14,7 62 bpd (29.5 percent ) of low

- octane C5+ naphtha. The diesel product can be sold as a commercial diesel fuel after the addition of appropriate additives. The C5+ naphtha product is a low octane material that is an excellent ethylene cracke r feedstock, or it can be further processed to make a gasoline blending component.

If more export power is required, additional coal must be consumed to produce it.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

52 Exhibit 4

- 1 Plant performance summary

– CO 2 sequestration case

Plant Output

Gas Turbine Power 111,600 kW
e Steam Turbine Power 315,600 kW
e

Total 427,200 kW
e Auxiliary Load Coal Handling 1,760 kW
e Coal Milling 8,160 kW
e Slag Handling 2,080 kW
e Air Separation Unit Auxiliaries 4,000 kW
e Air Separation Unit Main Air Compressor 235,480 kW
e

Oxygen Compressor 35,600 kW
e Nitrogen Compressors 6,080 kW
e CO 2 Compressor 43,480 kW
e Boiler Feedwater Pumps 1,640 kW
e Condensate Pump 200 kW
e Quench Water Pump 2,280 kW
e Syngas Recycle Compressor 2,920 kW
e Circulating Water Pump 6,800 kW
e Ground Water Pumps 800 kW
e Cooling Tower Fans 3,560 kW
e Scrubber Pumps 1,360 kW
e Acid Gas Removal 13,640 kW
e Gas Turbine Auxiliaries 4,000 kW
e Steam Turbine Auxiliaries 400 kW
e Claus Plant/TGTU Auxiliaries 1,000 kW
e Claus Plant TG Recycle Compressor 840 kW
e

F

- T Power requirement 31,948 kW
e Miscellaneous Balance of Plant A 12,000 kW
e

Transformer Losses 2,520 kW
e

Total 422,548 kW
e Plant Performance Net Auxiliary Load 422,548 kW
e Net Plant Power 4,652 kW
e Coal Feed Flowrate 794,022 (1,750,518) kg/hr (lb/hr) Thermal Input B 5,984,963 kW t Condenser Duty 928 (880) GJ/hr (MMBtu/hr)

Raw Water Withdrawal 32.8 (8,672) m 3 /min (gpm)

Carbon in coal 506,147 (1,115,863) kg/hr (lb/hr) Carbon in Slag 2,531 (5,579) kg/hr (lb/hr)

Sulfur 19,901 (43,875) kg/hr (lb/hr) Raw Water Consumption 26.8 (7,078) m 3 /min (gpm)

A

- Includes plant control systems, lighting, HVAC, and miscellaneous low voltage loads

B

- HHV of Illinois No. 6 coal is 27,135 kJ/kg (11,666 Btu/lb)Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

53 Exhibit 4

- 2 Plant performance summary

– CO 2 vent case

Plant Output

Gas Turbine Power 136,000 kW
e Steam Turbine Power 336,800 kW
e

Total 472,800 kW
e Auxiliary Load Coal Handling 1,760 kW
e Coal Milling 8,160 kW
e Slag Handling 2,080 kW
e Air Separation Unit Auxiliaries 4,000 kW
e Air Separation Unit Main Air Compressor 235,480 kW
e

Oxygen Compressor 35,600 kW
e Nitrogen Compressors 6,080 kW
e CO 2 Compressor 0 kW
e Boiler Feedwater Pumps 1,840 kW
e Condensate Pump 240 kW
e Quench Water Pump 2,280 kW
e Syngas Recycle Compressor 2,920 kW
e Circulating Water Pump 6,240 kW
e Ground Water Pumps 720 kW
e Cooling Tower Fans 3,280 kW
e Scrubber Pumps 1,360 kW
e Acid Gas Removal 13,640 kW
e Gas Turbine Auxiliaries 4,000 kW
e Steam Turbine Auxiliaries 400 kW
e Claus Plant/TGTU Auxiliaries 1,000 kW
e Claus Plant TG Recycle Compressor 840 kW
e

F

- T Power requirement 29,278 kW
e Miscellaneous Balance of Plant A 12,000 kW
e

Transformer Losses 2,520 kW
e

Total 375,718 kW
e Plant Performance Net Auxiliary Load 375,718 kW
e Net Plant Power 97,082 kW
e Coal Feed Flowrate 794,022 (1,750,518) kg/hr (lb/hr) Thermal Input B 5,984,963 kW t Condenser Duty 928 (880) GJ/hr (MMBtu/hr)

Raw Water Withdrawal 30.1 (7,940) m 3 /min (gpm)

Carbon in coal 506,147 (1,115,863) kg/hr (lb/hr) Carbon in Slag 2,531 (5,579) kg/hr (lb/hr)

Sulfur 19,901 (43,875) kg/hr (lb/hr) Raw Water Consumption 24.5 (6,473) m 3 /min (gpm)

A

- Includes plant control systems, lighting, HVAC, and miscellaneous low voltage loads

B

- HHV of Illinois No. 6 coal is 27,135 kJ/kg (11,666 Btu/lb)

4.1.1 Carbon Balance

An overall carbon balance for the facilities is provided in Exhibit 4

- 3 and Exhibit 4

- 4

. About

0.5 percent of the total carbon is encapsulated in the slag, while over 80 percent of the entering carbon is converted to syngas going to the F

- T reactor

. In the sequestration case t he compressed

CO 2 and F

- T liquid product streams contain the majority of carbon exiting the plant, accountingCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

54 for approximately 55 per cent and 40 percent of the total entering carbon, respectively. The total

amount of carbon lost to the environment is about 6 percent of the entering carbon.

Note that carbon emission s from the stack are in the form of CO 2 emissions to the atmosphere, and that these emissions only represent a fraction of the total life cycle GHG emissions

associated with the finished diesel fuel

.

Exhibit 4

- 3 Elemental carbon mass balance

– CO 2 sequestration case

Carbon In, kg/hr (lb/hr) Carbon Out, kg/hr (lb/hr)

Coal 506,109 (1,115,863) Naphtha 56,404 (124,349) 11%

Diesel 150,183 (331,096) 30%

Slag 2,531 (5,579) 1%

CO 2 272,389 (600,516) 54%

Stack gas 24,603 (54,324) 5%

Total 506,109 (1,115,863) Total 506,109 (1,115,863) 100 %

Exhibit 4

- 4 Elemental carbon mass balance

– CO 2 vent case

Carbon In, kg/hr (lb/hr) Carbon Out, kg/hr (lb/hr)

Coal 506,109 (1,115,863) Naphtha 56,404 (124,349) 11%

Diesel 150,183 (331,096) 30%

Slag 2,531 (5,579) 1%

Stack gas 297,030 (654,839) 59%

Total 506,109 (1,115,863) Total 506,109 (1,115,863) 100 %

4.1.2 Sulfur Balance

All the sulfur in the coal is recovered as elemental sulfur in the Claus plant.

4.1.3 Water Balance

Water balance s in Exhibit 4

- 5 and Exhibit 4

- 6 represent one of the four gasification section s (see

S ection 3.1.3 ) and one quarter of the F

- T are a water needs. The last r o w represents the complete

F

- T plant and all four of the gasification sections. The water demand is the total water demand for that plant section ; this consists of a combination of raw and recycled water. Some water may

also be discharged via a permitted discharge. The raw water consumption is the difference

between the water withdraw al and consumption. Raw water consumption is 8.4 and 7.7 barrels

of water per bbl of diesel produced for the CO 2 Sequestration Case and CO 2 Vent C ase, respectively ( 5.9 and 5.4 barrels of water per barrel of F

- T liquids produced).Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

55 Exhibit 4

- 5 Water b alance

– CO 2 sequestration case

Water Us age by Area

m 3 /min (gpm) Water

Demand Internal

Recycle Raw Water

Withdrawal Process

Water Discharge Raw Water

Consumption

Slag Handling 0.43 (114) 0.43 (114) 0.0 (0) 0.0 (0) 0.0 (0)

Quench/Wash 3.0 (786) 2.07 (548) 0.9 (238) 0.0 (0) 0.9 (238)

SWS Blowdown 0.0 (0) 0.0 (0) 0.0 (0) 0.03 (7)

- 0.03 (

- 7)

Condenser Makeup 0.8 (221) 0.0 (0) 0.8 (221) 0.0 (0) 0.8 (221)

Gasifier Steam 0.3 (84) 0.0 (0) 0.3 (84)

Shift Steam 0.5 (128) 0.0 (0) 0.5 (128)

BFW Makeup 0.03 ( 8 ) 0.0 (0) 0.03 ( 8 )

Cooling Tower 6.7 (1,7 42 ) 0.28 (7 4 ) 6. 3 (1,6 68 ) 1.5 (39 2 ) 4. 8 (1,2 76 )

BFW Blowdown 0 0.03 ( 8 )

- 0.03 (

- 8 )

SWS Blowdown 0.0 (0) 0.25 (66)

- 0.25 (

- 66)

SWS Excess Water 0.0 (0) 0.0 (0) 0.0 (0)

1⁄4 of F

- T ar ea values 0.16 (41.60) 0 0.16 (41.60) 0 0.16 (41.60)

Total 10.99 (2,9 04 ) 2.79 (73 6 ) 8.2 1 (2,1 68 ) 1.5 ( 398 ) 6.7 0 (1,7 70 )

Total 4 Gasifier sections plus F

- T plant 43.98 (11, 617 ) 11. 15 (2

, 94 5 ) 32.83 (8 ,672 ) 6 .03 (1 ,594 ) 26.79 (7 ,078 )

Exhibit 4

- 6 Water b alance

– CO 2 vent case

Water Usage by Area

m 3 /min (gpm) Water

Demand Internal

Recycle Raw Water

Withdrawal Process

Water Discharge Raw Water

Consumption

Slag Handling 0.43 (114) 0.43 (114) 0.0 (0) 0.0 (0) 0.0 (0)

Quench/Wash 3.0 (786) 2.07 (548) 0.9 (238) 0.0 (0) 0.9 (238)

SWS Blowdown 0.0 (0) 0.0 (0) 0.0 (0) 0.03 (7)

- 0.03 (

- 7)

Condenser Makeup 0.8 (222) 0.0 (0) 0.8 (222) 0.0 (0) 0.8 (222)

Gasifier Steam 0.3 (84) 0.0 (0) 0.3 (84)

Shift Steam 0.5 (128) 0.0 (0) 0.5 (128)

BFW Makeup 0.03 (9) 0.0 (0) 0.03 (9)

Cooling Tower 6.1 (1,601) 0.28 (75) 5.8 (1,526) 1.4 (360) 4.4 (1,166)

BFW Blowdown 0.0 (0) 0.03 (9)

- 0.03 (

- 9)

SWS Blowdown 0.0 (0) 0.25 (66)

- 0.25 (

- 66)

SWS Excess Water 0.0 (0) 0.0 (0) 0.0 (0)

1⁄4 of F

- T area values 0.04 (9) 0.04 (9) 0.0 (0) 0.0 (00) 0.0 (00)

Total 10.3 (2,722) 2.79 (737) 7.5 (1,985) 1.4 (367) 6.1 (1,618)

Total 4 Gasifier sections plus F

- T plant 41.3 (10,896) 11.20 (2956) 30.1 (7,940) 5.6 (1,467) 24.5 (6,473)Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

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56 

## 5 C ost Estimation Results

5.1 Capital and Operating Costs

The gasifier, associated syngas treatment units, Rectisol ® system, CO 2 compressors, Claus unit,

g as and steam turbines, and other equipment typically associated with an IGCC plant were scaled from updated cost estimates (2011 dollar basis) for the Shell gasifier C ase 6 in “ Cost and

Performance Baseline for Fossil Energy Power Plants, Volume 1: Bituminous Coal and Natural Gas to Electricity

. ” [ 8 ] Capital costs for the F

- T reactor and various treatment and upgrading

systems were based on adjusted, vendor

- furnished, actual cost data, and best possible projections.

A process contingency of 25 percent was applied to the F

- T reactor and associated equipment. The capital costs at the TPC level include equipment, materials, labor , indirect construction costs, engineering, and contingencies. Additional details are highlighted in S ection 1.6 of this report and the NETL QGESS. [ 6 ]

The TOC for the entire CTL facility with CO 2 capture was estimated to be $ 6,53 2 million . This

cost includes appropriate contingency factors and owners costs

. [ 6 ] On a per barrel basis of daily liquid capacity, this cost is $ 130, 662 / bpd FTL for the sequestration case

. Exhibit 5

- 1 shows a summary breakdown of the capital costs for the CO 2 sequestration case and Exhibit 5

- 3 shows a summary breakdown of the capital costs for the CO 2 vent case . The total TOC is the TPC plus

owner’s costs.

A credit of $ 60 /MWh is included for excess power generated. Exhibit 5

- 2 and Exhibit 5

- 4 show the breakdown of the total operating cost for the sequestration and vent cases, respectively

.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

57 Exhibit 5

- 1 Total plant cost summary with CO 2 s equestrationCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

58 Exhibit 5

- 1 Total plant cost summary with CO 2 sequestration (continued )Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

59 Exhibit 5

- 1 Total plant cost summary with CO 2 sequestration (continued )Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

60 Exhibit 5

- 1 Total plant cost summary with CO 2 sequestr ation (continued )Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

61 Exhibit 5

- 1 Total plant cost summary with CO 2 sequestration (continued )Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

62 Exhibit 5

- 2 Total operating and maintenance costs with CO 2 s equestrationCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

63 Exhibit 5

- 3 Total plant cost summary

– CO 2 v entCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

64 Exhibit 5

- 3 Total plant cost summary

– CO 2 vent (continued )Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

65 Exhibit 5

- 3 Total plant cost summary

– CO 2 vent (continued )Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

66 Exhibit 5

- 3 Total plant cost summary

– CO 2 vent (continued )Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

67 Exhibit 5

- 3 Total plant cost summary

– CO 2 vent (continued )Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

68 Exhibit 5

- 4 Total operating and maintenance costs

– CO 2 v entCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

69 

### 5.2 Cost of Pr oduction

Two financial structures were examined that reflect a hypothetical future in which technical risk has been partially mitigated by the demonstration of two or three commercial

- scale CTL plants. The first is assumed to have no government incentives (commercial fuels) while moderate government incentives

– in the form of a government loan guarantee

– are assumed to be available under the second. Both scenarios face a high degree of market risk because of the volatile nature of energy prices, especia lly the world oil price. The TOC, TASC, and annual

O&M costs were estimated based on plant performance and a 90 percent capacity factor (CF). The COP under each financial scenario structure were calculated. The results are presented in

Exhibit 5

- 5

. These costs include a credit of $ 60 /MWh for excess power generated. These values are consistent with those expected in an environment after a carbon regulatory framewor k has

been established. T he electricity credit accounts for less than 0.5 percent of the total COP in the sequestration case and less than 2.5 percent of the total COP in the vent case, as shown in

Exhibit 5

- 6

, any change in the assumed price would have very little impact on the overall COP

values calculated.

Exhibit 5

- 7 provides the TOC per F

- T bbl, ECO bbl, and EPD bbl values for each case for ease of comparison.

Exhibit 5

- 5 Cost estimation results summary ( CO 2 s equestration )

CO 2 Venting Case CO 2 Sequestration Case

Commercial

Fuels Loan

Guarantees Commercial

Fuels Loan

Guarantees

Total Overnight Cost, $1000 6,398,760 6,532,052
Total Overnight Cost, $/bpd 127,996 130,662
Total Overnight Cost, $/bpd ECO 122,014 124,557
Total Overnight Cost, $/bpd EPD 152,518 155,696

Total As Spent Cost, $1000 7,558,216 7,341,938 7,715,660 7,494,877

Total As Spent Cost, $/bpd 151,189 146,862 154,338 149,922

Total As Spent Cost, $/bpd ECO 144,123 139,999 147,127 142,917

Total As Spent Cost, $/bpdl EPD 180,154 174,999 183,908 178,646

Total annual fixed O&M, $1000 197,405 197,405 200,854 200,854

Total annual variable O&M (90% CF), $1000 127,370 127,370 130,076 130,076

Total annual feedstock cost (90% CF), $1000 473,377 473,377 473,377 473,377

Total annual power credit (90% CF), $1000

- 44,844

- 44,844

- 2,149

- 2,149

COP F

- T Diesel, $/bbl FTD 143.7 123.1 154.2 133.2

COP F

- T Naphtha, $/bbl FTN 100.1 85.8 107.4 92.8

Equivalent Crude Oil Price, $/bbl ECO 124.7 106.9 133.8 115.6

Equivalent Petroleum Diesel Price, $/bbl EPD 155.9 133.6 167.2 144.5

All costs in June 2011 dollars.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

70 Exhibit 5

- 6 COP sensitivity to market price of electricityCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

71 Exhibit 5

- 7 Comparison of TOC per F

- T bbl, ECO bbl, and EPD bbl

Source: NETLCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

72 Exhibit 5

- 8 provides a chart that demonstrates the sensitivity of t he cost of production of FT

Liquids to the capital charge factor. The capital charge factors (CCF) correlate with various finance structures, as pointed out in the chart, with high

- risk investor owned utility ( IOU ) having a CCF of 0.124, loan guarantees h aving a CCF of 0.170, commercial fuels having a CCF of

0.218 and 100 percent equity at 20 percent internal rate of return on equity ( IRROE ) having a

CCF of 0.350.

Exhibit 5

- 8 FT

- Liquids COP sensitivity to CCF

Source: NETLCost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

73 

### 5.3 Cost of Carbon Capture

The base case F

- T process with carbon sequestration was compared to the case w here CO 2 was vented

. In the vent case t he partial shift and carbon separation f rom syngas using Rectis ol ® was kept to maintain the H 2 /CO ratio fed to F

- T reactor as was the amine

- based tail gas CO 2 capture system to prevent the build

- up of inerts in the recycle loop

. However, t he WGS reactors and

amine

- based capture process were removed from the fuel gas stream as was the CO 2 compressors, pipeline, and cost of sequestration. Exhibit 5

- 9 summa rizes the additional costs due to carbon capture.

The cost of capture was calculating using the following equation:

Cost of Carbon Capture = ( A

− B ) ( C

D ) ⁄

Where: A = C ost of production for Fischer Tropsch Diesel with sequestration ( $/bbl FTD )

B = C ost of production for Fischer Tropsch Diesel without sequestration ( $/bbl FTD )
C = Rate of CO 2 sequestration (tonne/day) D = Sum of total diesel daily production and adjusted total naphtha production. Naphtha is adjusted by a factor of 0.70 to account f or lower value.

Exhibit 5

- 9 Cost summary of CTL f acility with and without CCS

Vent Case Sequestration

Case Cost of adding

CCS (delta)

Total Overnight Cost, $1000 6,398,760 6,532,052 133,292

Total annual fixed O&M, $1000 197,405 200,854 3,448

Total annual variable O&M (90% CF), $1000 127,370 130,076 2,706

Total annual fuel cost (90% CF), $1000 473,377 473,377 0

COP F

- T Diesel * , $/bbl FTD 123.1 [143.7] 133.2 [154.2] 10 .1 [10.5]

COP F

- T Naphtha *

, $/bbl FTN 85.8 [100.1] 92.8 [107.4] 7 .0 [7.3]

Equivalent Crude Oil Price * , $/bbl ECO 106.9 [124.7] 115.6 [133.8] 8.7 [9.1]

Equivalent Petroleum Diesel Price * , $/bbl EPD 133.6 [155.9] 144.5 [167.2] 10.9 [ 11 .3]

Cost of Carbon Capture*, $/Tonne N/A 19.11 [19.94] 19.11 [19.94]

* Values are shown for two financial structures. The first (lower) value is based on the loan guarantee finance structure. The second (higher/bracketed) value is based on the commercial fuels finance

structure.
All costs in June 201 1 dollars.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

74 

## 6 Conclusions and Recommendations

6.1 Conclusions

Developing an alternative source of liquid fuels using coal as the carbon input has national security implications and economic consequences. The results of this study show that liquid fuels produced from a 50,000 bbl/day coal

- based F

- T plant (low tempera ture, iron

- based catalyst) are more expensive than fuels produced conventionally from crude oil at today’s oil prices (about $98/barrel, WTI average in 2013). Government incentives in the form of loan guarantees significantly reduce the cost of F

- T liquid s, but not to less than current fuel prices derived from crude oil. The two cases examined in this study, one with CO 2 captured and one

with CO 2 vented, have ECO prices of $134/bbl and $125/bbl, respectively, using the commercial fuels financing structure. Government loan guarantees reduce the ECO price for the same two cases to $116/bbl and $107/bbl, respectively. Barring a dramatic shift in the oil market, and given the extremely large capital investment required for these CTL plants (on the order of $6 .5 billion TOC), loan guarantees will likely be required to incentivize project development.

CO 2 capture is an inherent part of the F

- T process resulting in two point sources: from the shifted syngas via the Rectisol process and from the F

- T tail gas via a n amine

- based process. In the vent case, CO 2 from these two sources was released to the atmosphere. In the sequestration case

, CO 2 from these two sources was compressed to 2,214 psia along with additional CO 2 captured from

the combined cycle power s ectio n . The incremental cost of adding this capture and compression

for sequestration is about $9/bbl (ECO price) independent of the financing structure used.



### 6.2 Recommendations for Further Study

Suggestions for future work include :

 Evaluate plant performance and economics using different coal types (subbituminous and lignite).

 Perform a detailed water usage analysis of the facility to determine more accurately the water makeup and discharge rates. Also determine the best means for minimizing water usage.

 Perfor m a literature search to identify ways of reducing nitrogen and a rgon concentration in

s yngas to reduce recycle load to the F

- T reactor.Cost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

75 

## 7 References

1 “Baseline Technical and Economic Assessment of a Commercial Scale Fischer

- Tropsch Liquids Facility,” DOE/NETL

- 2007/1260, Final Report for Subtask 41 817.401.08.001, National Energy Technology Laboratory

. April 2007. Available at: http://netl.doe.gov/research/energy

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DOE/NETL

- 2009/1349, National Energy Technology Laboratory, Pittsburgh, PA. January 2009. Available at: http://www.netl.doe.gov/File%20Library/Research/Energy%20Analysis/Coal/CBTL

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6 National Energy Technol ogy Laboratory (NETL). (2011). QGESS: Cost Estimation Methodology for NETL Assessments of Power Plant Performance . Energy Analysis

Publications: Details. Retrieved on March 16, 2012 , from http://www.netl.doe.gov/research/energy

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guidelines

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8 NETL. “ Cost and Performance Baseline for Fossil Energy Plants, Volume 1: Bituminous Coal and Natural Gas to Electricity”. DOE/NETL

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, Final Report

. November 2010

Available at: http://www.netl.doe.gov/File%20Library/Research/Energy%20Analysis/OE/BitBase_Fin Rep_Rev2a

- 3_20130919_1.pdf

9 AACE (formerly Advancement of Cost Engineering International ) (2005) Cost Estimate Classification System

– As Applied in Engineering, Procurement, and Construction for the Process Industries; TCM Framework 7.3

– Cost Estimating and Budgeting , AACE

International Recommended Practice No. 18R

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Economic Evaluations

– as Applied for the Process and Utility Ind u stries TCM

Framework: 3.2

– Asset Planning, 3.3

– Investment Decision Making , AACECost and Performance Baseline for Fossil Energy Plants Volume 4: CTL via F

- T Synthesis

76 International Recommended Practice No. 16R

- 90, 2003. Retrieved on April 8, 2013 from http://www.aacei.org/non/rps/16R

- 90.pdf

11 NETL Power Systems Financial Model Version 5.0, December 2008. User Guide available at http://www.netl.doe.gov/research/energy

- analysis/publications/list ing?progcat=all

12 Wimer, J. Recommended Project Finance Structures for the Economic Analysis of Fossil

Based Energy Projects. September 2008.

13 Tarka, T. Polygeneration of Fischer

- Tropsch Diesel Fuel and Electricity from Coal U.S. Department of Energy, National Energy Technology Laboratory. August 2010.

14 Stefan Rabe, Thanh

- Binh, Truong, et al. Low Temperature Catalytic Partial Oxidation of Methane for Gas

- to

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15 A STM

, TABLE X5.1 Tenth Percentile Minimum Ambient Air Temperatures for the United States (except Hawaii)

, American National Standard Designation: D 975

– 07 , 2007

. 16 Worldwide Fuel Charter Committee

, World Wide Fuel Charter, Fifth Edition , 2013

.