EPA- 600/7- 90-018;
September 1990
ASSESSMENT OF CONTROL TECHNOLOGIES FOR REDUCING
IONS OF S02 AND NOX FROM EX1
COAL-FIRED UTILITY BOILERS
EMISSIONS OF S02 AND NOX FROM EXISTING
FINAL REPORT
by
David M. White and Mehdi Maibod1
Radian Corporation
Post Office Box "13000
Research Triangle Park, North Carolina 27709
EPA Contract No. 68-02-4286
Work Assignment 84
Project Officer
Norman Kaplan
U. S. Environmental Protection Agency
Air and Energy Engineering Research Laboratory
Research Triangle Park, North Carolina 27711
This Project Was Conducted in Association With the
National Acid Precipitation Assessment Program
Prepared for
U. S. Environmental Protection Agency
Office of Research and Development
Washington, O.C. 20460
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TECHNICAL REPORT DATA
(Please read Instructions on the reverse before completi"
I. REPORT NO.
EPA-600/7-90-018
2.
PB90-27357U
v
4. TITLE AND SUBTITLE
Assessment of Control Technologies for Reducing
Emissions of SO2 and NOX from Existing Coal-fired
Utility Boilers
7. AUTHOR(S)
David M. White and Mehdi Maibodi
5. REPORT DATE
September 1990
6. PERFORMING ORGANIZATION CODE
8. PERFORMING ORGANIZATION REPORT NO
9. PERFORMING ORGANIZATION NAME AND ADDRESS
Radian Corporation
P. O. Box 13000
Research Triangle Park, North Carolina 27709
10. PROGRAM ELEMENT NO.
11. CONTRACT/GRANT NO.
68-02-4286, Task 84
12. SPONSORING AGENCY NAME AND ADDRESS
EPA. Office of Research and Development
Air and Energy Engineering Research Laboratory
Research Triangle Park, North Carolina 27711
13. TYPE OF REPORT AND PERIOD COVERED
Task final; 1/87 - 12/89
14. SPONSORING AGENCY CODE
EPA/600/13
is.SUPPLEMENTARY NOTES AEERL project officer is Norman Kaplan, Mail Drop 62, 919/541-
f\ r- ^ rt
2556.
\
16. ABSTRACT-Vr,, . ,n . , . . , . . .. _ :
The report reviews available information and estimated costs on 15 emission
control technology categories applicable to existing coal-fired electric utility boilers.
The categories include passive controls such as least emission dispatching, conven-
tional processes, and emerging technologies still undergoing pilot scale and commer-
cial demonstration. The status of each technology i.s-reviewed relative to four ele-
ments;-'(l)£Description--how the technology works;-*(2^Applicability--its applicability
to existing plants;J(3^Performance--the expected emissions reduction; and-j(4)^:Costs
--the capital cost, busbar cost, and cost per ton of SO2 and NOx removed. Costs
are estimated for new and retrofit applications for various boiler sizes, operating
^Characteristics, fuel qualities, and boiler retrofit difficulties. Capital costs vary
from($2/kW for overfire air to(j£2800/kW for integrated gasification combined cycle
in 1988 dollars.^NOTE: A major objective of the National Acid Precipitation Asses-
sment Program-- is to evaluate alternative methods for reducing SO2 and NOx emis-
sions from combustion sources and to identify options which appear most promising
from both an emissions reduction and cost standpoint. Part of this overall effort is
to develop up-to-date generic assessments of commercial, near-commercial and
emerging emission control technology categories applicable to these utility boilers.)
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b.lDENTIFIERS/OPEN ENDED TERMS
c. COSATl Field/Group
Pollution
Utilities
Boilers
Coal
Combustion
Sulfur Dioxide
Nitrogen Oxides
Cost Effectiveness
Gasification
Emission
Pollution Control
Stationary Sources
13 B
13A
21D
21B
07B
.14A
14H, 07A
14G
8. DISTRIBUTION STATEMENT
Release to Public
19. SECURITY CLASS (ThisReport)
Unclassified
21. NO. OF PAGES
20. SECURITY CLASS (Thispage)
Unclassified
22. PRICE
EPA Form 2220-1 (9-73)
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NOTICE
This document has been reviewed in accordance with
U.S. Environmental Protection Agency policy and
approved for publication. Mention of trade names
or commercial products does not constitute endorse-
ment or recommendation for use.
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ABSTRACT
A major objective of the National Acid Precipitation Assessment Program
is to evaluate alternative methods for reducing S(L and NO emissions from
combustion sources and to identify those options which appear most promising
from both an emissions reduction and cost standpoint. Part of this overall
effort is to develop up-to-date generic assessments of commercial, near-
commercial, and emerging emission control technologies applicable to existing
coal-fired electric utility boilers. This report reviews available information
and estimated costs on 15 technology categories, including passive controls
such as least emission dispatching, conventional processes, and emerging
technologies still undergoing pilot scale and commerical demonstration.
The status of each technology is reviewed relative to the following four
elements:
t Description—how does the technology work?
• Applicability—what is its applicability to existing plants?
• Performance--what is the expected emissions reduction?
t Cost—what is the capital cost, busbar cost, and cost
per ton of StL and NO removed?
Cost estimates are presented for new and retrofit applications for various
boiler sizes, operating characteristics, fuel qualities, and boiler retrofit
difficulty. Capital costs vary from $2 per kilowatt for Overfire Air to $2,800
per kilowatt for Integrated Gasification Combined Cycle in 1988 dollars.
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CONTENTS
Abstract ;... j.'-;. . iii
Figures -....' :X . vi
Tabl es !.<£ . y.f.'.. i x
Abbrevi ations and Symbol s y.'... ;•->.. xi
1. Introduction and Summary 1-1
References 1-12
2. Commercial Technologies 2-1
2.1 Fuel Switching and Blending 2-2
2.2 Least Emissions Dispatch 2-15
2.3 Physical Coal Cleaning 2-18
2.4 NO Combustion Control Technology 2-28
2.5 Thfowaway Wet FGD 2-36
2.6 By-Product Recovery FGD Technologies 2-47
2.7 Spray Drying 2-57
References '. 2-68
3. Near-Commercial Technologies 3-1
3.1 Integrated Gasification Combined Cycle 3-2
3.2 Fluidized Bed Combustion 3-13
3.3 Post-Combustion NO Control... 3-25
3.4 Furnace Sorbent Injection 3-33
3.5 Low-Temperature Sorbent Injection 3-44
3.6 Reburning 3-51
References 3-56
4. Emerging Technologies .4-1
4.1 Advanced Coal Cleaning 4-2
4.2 Advanced Post-Combustion S0?/N0 Processes..... 4-8
References... T. 4-16
Appendices
A. Summary of Control Costs Coal Switching
and Blending A-l
B. Low NO Combustion Control Technologies B-l
C. Lime/Limestone FGD C-l
D. Integrated Gasification Combined Cycle • D-l
E. Atmospheric Fluidized Bed E-l
F. Lime Spray Drying — F-l
G. Selective Catalytic Reduction G-l
H. Furnace Sorbent Injection. H-l
I. Natural Gas Reburn 1-1
Preceding, page Wank
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FIGURES
Figure Page
1-1 Capital Costs - Constant 1988 Dollars 1-9
1-2 Levelized Annual Cost - Constant 1988 Dollars 1-10
1-3 Unit Cost - Constant 1988 Dollars 1-11
2.1-1 Effect of Coal Rank on Furnace Sizing 2-3
2.1-2 Coal Supply Regions 2-7
2.1-3 Coal Switching - Capital Cost 2-12
2.1-4 Coal Switching - Levelized Annual Cost 2-12
2.1-5 Coal Switching - Cost Per Ton of S02 Removed 2-13
2.3-1 Coal Washability Curves 2-20
2.3-2 Level 4 Coal Preparation Plant 2-23
2.4-1 NOX Combustion Controls - Capital Cost.. 2-34
2.4-2 NO Combustion Controls - Levelized Annual Cost 2-34
2.4-3 NOX Combustion Controls - Cost Per Ton of NOX Removed 2-35
2.5-1 Lime/Limestone FGD System Flow Diagram 2-37
2.5-2 Simplified Flow Diagram for a Dual Alkali
FGD System 2-39
2.5-3 L/LS FGD - Capital Cost 2-45
2.5-4 L/LS FGD - Levelized Annual Cost 2-45
2.5-5 L/LS FGD - Cost Per Ton of S02 Removed 2-46
2.6-1 Wellman-Lord Process Schematic ; 2-48
2.6-2 Magnesia Slurry Process Schematic 2-50
2.7-1 Lime Spray Drying Process Flow Diagram 2-58
2.7-2 Lime Spray Drying - Capital Cost 2-64
VI i
—i
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FIGURES
(Continued)
jure Page
2.7-3 Lime Spray Drying - Annualized Cost 2-65
2.7-4 Lime Spray Drying - Cost Per Ton of S02 Removed 2-66
3.1-1 Generalized Block Flow Diagram of Combined Cycle Coal
Gasification Power Generation 3-3
3.1-2 IGCC - Capital Cost 3-11
3.1-3 IGCC - Level ized Annual Cost 3-11
3.2-1 Simplified AFBC Process Flow Diagram 3-14
3.2-2 AFBC - Capital Cost 3-23
3.2-3 AFBC - Level ized Annual Cost 3-23
3.3-1 Possible SCR Configurations 3-26
3.3-2 Selective Catalytic Reduction - Capital Cost.. 3-31
3.3-3. Selective Catalytic Reduction - Levelized Annual Cost 3-31
3.3-4 Selective Catalytic Reduction - Cost Per Ton of NOX Removed... 3-32
3.4-1. Simplified Schematic of Furnace Sorbent Injection 3-34
3.4-2 Peak Sorbent Reactivity as a Function Of Temperature 3-34
3.4-3 S02 Removal as a Function of Calcine Surface Area 3-37
3.4-4 S02 Removal as a Function of Ca/S Ratio and Residence Time 3-37
3.4-5 Increase in Solids Loading as a Function Of Coal Sulfur
Content and Ca/S Ratio for a Typical 10% Ash Coal 3-40
3.4-6 Furnace Sorbent Injection - Capital Cost 3-42
3.4-7 Furnace Sorbent Injection - Level ized Annual Cost 3-42
3.4-8 Furnace Sorbent Injection - Cost Per Ton of S02 Removed 3-43
3.5-1 SO, Removal as a Function of Normalized Stoichiometric Ratio
(NSR) 3-45
vii
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FIGURES
(Continued)
Figure Page
3.6-1 Natural Gas Reburning - Capital Cost 3-54
3.6-2 Natural Gas Reburning - Levelized Annual Cost 3-54
3.6-3 Natural Gas Reburning - Cost Per Ton of NOX Removed 3-55
4.2-1 E-Beam/Ammonia Process Flow Diagram 4-9
4.2-2 Simplified Flow Diagram for the Fluidized-Bed Copper
Oxide Process 4-11
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TABLES
Table Page
1-1 Control Technologies Reviewed 1-2
1-2 Bases for Cost Estimates 1-4
1-3 Summary of Cost Resul ts 1-8
2.1-1 Technical Factors Affecting Coal Switching 2-4
2.1-2 Typical Delivered Coal Properties for Selected
Supply Regions 2-8
2.1-3 Factors Influencing the Decision to Switch or
Scrub to Reduce Sulfur Dioxide Emissions 2-14
2.3-1 Breakdown of Coal Cleaning Plants by Region
and Level, 1978 2-24
2.3-2 Potential Reductions in Sulfur Dioxide Emission Rates from
Physical Coal Cleaning 2-24
2.3-3 Economics of Conventional Physical Coal Cleaning 2-27
2.4-1 Site-Specific Parameters Affecting NO Emissions from
Combustion Modification Controls 2-30
2.5-1 Total Operating and Planned Throwaway FGD Capacity
by U. S. Electric Utilities (as of December 1985) 2-42
2.6-1 Wellman-Lord Utility Installations in the U. S..... 2-52
2.6-2 Cost Estimates for Wellman-Lord FGD.. 2-55
3.1-1 Available S0~ Emissions Data for Cool Water
Demonstration Program 3-6
3.1-2 Design Sulfur Emissions Information for EPRI IGCC
Studies 3-8
3.1-3 Available NO Emissions Data for Cool Water
Demonstration Program 3-9
3.1-4 Summary of Available Emission and Cost Data for IGCC
Facilities 3-10
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TABLES
(Continued)
Table Page
3.2-1 Estimated Operating Data for 200 MW AFBC 3-17
3.2-2 Full-Scale Utility FBC Demonstrations : 3-18
3.2-3 Comparative Economics of New 500 MW Conventional and FBC
Power Plants3.' 3-22
3.4-1 SO, Capture as a Function of Ca/S Ratio, Quench Rate, and
Sofbent 3-36
4.1-1 Economics of Advanced Physical Coal Cleaning 4-7
4.1-2 Economics of Chemical Coal Cleaning 4-7
4.2-1 Cost Estimates for Electron Beam 4-14
4.2-2 Cost Estimates for Copper Oxide 4-15
i Y I
x
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ABBREVIATIONS
ABBREVIATIONS AND SYMBOLS
ADVACATE
AEERL
AFBC
AFDC
ASTM
AUSM
CCR
CCTF
CF
CS/B +$5
CS/B +$15
DM
DOE
DSD
EER
EPRI
ESP
FBC
FF
FPD
FSI
GRI
HALT
HRSG
IAPCS
IGCC
kW
L/LS FGD
LED
LNC
LNB
LTSI
LSD
MHI
MW
NAPAP
NESCAUM
NGR
NSPS
advanced silicate process
Air and Energy Engineering Research Laboratory
atmospheric fluidized bed combustion
allowance for funds during construction
American Society for Testing and Materials
advanced utility simulation model
Conoco Coal Research
coal cleaning test facility
capacity factor
coal switching and blending at $5 coal cost
differential
coal switching and blending at $15 coal cost
differential
dense media
(U.S.) Department of Energy
duct spray drying
Energy and Environmental Research Corporation
Electric Power Research Institute
electrostatic precipitator
fluidized bed combustion
new fabric filter
fuel price differential
furnace sorbent injection
Gas Research Institute
hydrate addition at low temperature
heat recovery steam generator
integrated air pollution control system
integrated gasification combined cycle
kilowatt
lime/limestone flue gas desulfurization
least emissions dispatch
low NOx combustion
low NOx burners
low-temperature sorbent injection
lime spray drying
Mitsubishi Heavy Industries
megawatt
National Acid Precipitation Assessment Program
Northeast States for Coordinated Air Use Management
natural gas reburn
new source performance standards
xi
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ABBREVIATIONS
(Continued)
NSR -- normalized sto1chiometr1c ratio
OFA -- overfire air
PCC -- physical coal cleaning
PM -- particulate matter
PFBC -- pressurized fluldized bed combustion
SCA -- specific collection area
SCR -- selective catalytic reduction
S.G. -- specific gravity
STAR -- state acid rain
U.S. EPA -- U.S. Environmental Protection Agency
CONVERSION TO METRIC VALUES
Btu -- 1,054.8 Joules
Btu/lb -- 2,325.445 Joules/Kg
"3'
ft2/1000 acfm -- 0.3048 m2/actual m3/sec
gallon -- 0.0038 m
pound (Ib) -- 0.45359 Kg
lb/106Btu -- 0.0043 Kg/Joules
ton -- 907.18 Kg
xii
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INTRODUCTION AND SUMMARY
BACKGROUND AND PURPOSE
One of the objectives of the National Acid Precipitation Assessment
Program is to evaluate the potential performance and cost of alternative
methods for reducing S02 and NO emissions from combustion sources. Part of
this overall effort is to develop up-to-date generic information on commer-
cial, near-commercial, and emerging emission control technologies applicable
to coal-fired electric utility boilers. This report presents a review of
available information on the technologies shown on Table 1-1. Because the
various acid rain regulatory proposals focus on reduction of S0? and NO in
Cf rt
the eastern half of the United States, the report focuses primarily on each
technology's potential for retrofit onto existing boilers in the eastern
U. S. burning medium and high sulfur coals.
ORGANIZATION
The technology reviews are divided into three major sections covering
technologies which are commercial (Section 2), near-commercial (Section 3),
and emerging (Section 4). These three classes are respectively defined, as
follows: technologies routinely used by U. S. electric utilities,
technologies undergoing large-scale demonstration by U. S. utilities or
commercially used in Japan or Europe, and those still undergoing laboratory
or pilot-scale testing. Designation of a technology to one of these three
classes is based on the technology's demonstrated status on low and high
sulfur coals.
Within each major section, the technologies are presented in the
following order: passive controls, pre-combustion controls, combustion
controls, post-combustion controls, and combined systems. The term "passive
controls" refers to technologies which in many cases require little or no
1-1
-------
TABLE 1-1. CONTROL TECHNOLOGIES REVIEWED
Section
Technology
Potential Emission
Reductions (%)
SO,
NO.
Commercial :
2.1
2.2
2.3
2.4
2.4
2.5
2.5
2.5
2.6
2.7
Near Commercial :
3.1
3.2
3.3
3.4
3.5
3.6
Emerging;
4.1
4.2
4.2
Fuel Switching and Blending
Least Emissions Dispatch
Physical Coal Cleaning
Low NO Burners
OverfiPe Air
Lime/Limestone FGO
Additive Enhanced FGD
Dual Alkali FGD
By-Product Recovery FGD
Spray Drying
Integrated Gasification
Combined Cycle
Fluidized Bed Combustion
Selective Catalytic Reduction
Furnace Sorbent Injection
Low-Temperature Sorbent Injection
Reburning
Advanced Coal Cleaning
Electron Beam Irradiation
Copper Oxide FGD
50-80
0-90
20-50
0
0
90-95
90-95
90-95
90-95
70-90
90-95
80-90
0
50-70
50-70
15-20
45-60
80-95
90-95
0-10
0-40
0
30-50
15-30
0
0
0
0
0
90-95
>50
80-90
0
0
35-50
0
55-90
90-95
FGD - Flue Gas Desulfurization.
1-2
-------
capital expenditure (i.e., hardware), but which will require changes in a
utility's operating methods. The status of each technology is reviewed
relative to the following four elements:
• Description—how does the technology work?
• Applicability--what is its applicability to existing plants
burning low and high-sulfur coals?
• Performance—what is the expected emissions reduction?
• Cost—what are the capital cost, busbar cost, and cost
per ton of S0? and NO removed?
£ A
Because of the importance of consistent treatment of each technology, a
consistent set of economic procedures was used for most technologies to
allow comparisons. The methodology used for this purpose is discussed
below.
METHODOLOGY
Because of the diversity of plant sizes and designs, operating charac-
teristics, fuel quality, and financing arrangements found throughout U. S.
utilities, it was necessary to define a uniform methodology for use in the
report. These procedures can be divided into two major categories: boiler
design and economic assumptions. Base case, high, and low values were
selected for boiler design and economic parameters. The range in values was
evaluated to present boiler conditions which may favor the selection of one
technology option over another. Table 1-2 presents the range of boiler
design and economic assumptions selected.
For the technologies addressed in this report, order of magnitude cost
estimates are presented. Cost estimates presented in the text are based on
a range of boiler and coal parameters. Cost of many technologies is very
site-specific and varies significantly depending on the boiler and coal
characteristics.
The Integrated Air Pollution Control Systems (IAPCS) (1) cost model,
which is currently being updated to include more technologies, was used to
develop the cost estimates for some of the technologies in this report.
1-3
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TABLE 1-2. BASES FOR COST ESTIMATES3
Parameter Descriptions
Base Case
Value
High Case
Value
Low Case
Value
Boiler and Coal Characteristic Assumptions
Unit:
Size in MM 300 100 700
Capacity factor, % 50 10 90
Specific collection area 300 200. 400
Coal Characteristics:
Sulfur content, % 2.0 1.0 4.0
Switched fuel sulfur content, % 0.9 0.9 0.9
Ash content, % 10.0 5.0 15.0
High heating value, Btu/lb 11,000 9,000 13,000
Economic Assumptions
Capital Cost Indirects:
General facilities, % 10%
Engineering, % 10%
Project contingencies, % 30%
Process contingencies, % 0-10%, commercial technologies
10-30%, developing technologies
Retrofit factor (for FGD or SCR) 1.3 1.5 1.0
Economic Life ... 20 15 30
Carrying Charge Factor 0.189 0.205 0.175
O&M Levelizing Factor 1.57 1.45 1.75
Operating Costs:
Fuel price differential,
Operating labor, $/hr
Steam, S/1000 Ib
Electricity, mills/kWh
Lime, $/ton
Limestone, $/ton
Organic acid, $/ton
Ammonia, $/ton
SCR catalyst, $/ton
Waste disposal, $/ton
Water, 5/1000 gal.
Natural gas, $/10° Btu
Sulfur, $/ton
$/ton 10 15 5
21.4
7.0
57.0
60.0
16.0
1,725.0
150.0
20,300.0
10.0
0.65
2.0
65.0
It is EPA's policy to use metric units. However, for the convenience of
the reader non-metric units are used in this report. Conversion factors
to metric units are given in the table of Abbreviations and Symbols.
1-4
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The cost/performance assumptions are the same as used under the National
Acid Precipitation Assessment Program site-specific retrofit cost study
under which the costs of retrofitting SO- and NO controls at 200 coal-fired
utility power plants are being estimated (2). The IAPCS cost model was used
to develop cost estimates for the following control technologies.
• Coal switching and blending (CS/B),
• Furnace sorbent injection (FSI),
t Lime spray drying with reuse of the existing electrostatic
precipitators (LSD+ESP),
• Lime spray drying with a new fabric filter (LSD+FF),
• Lime/limestone FGD (L/LS FGD),
t Natural gas reburn (NGR),
t Low NO burner (LNB),
• Overfire air (OFA),
t Selective catalytic reduction (SCR),
• Integrated gasification combined cycle (IGCC), and
t Atmospheric fluidized bed combustion (AFBC).
For the other technologies addressed in this report, costs are from
referenced publications. These costs are not included in this section for
comparison, since other cost model assumptions were used in generating costs
which may not be consistent with assumptions used 1n the IAPCS cost
estimates.
Economic Assumptions
Cost estimates are presented in 1988 dollars using both current and
constant dollar procedures. The Electric Power Research Institute's (EPRI)
general costing procedures were used to incorporate inflation, cost of
capital, and levelization of future expenses (3). The cost of replacement
power or lost capacity while a plant is out-of-service during retrofit is
not included in the analysis. Downtime replacement power costs depend on
the duration of the downtime period and the cost differential between the
purchased or replaced electricity and the cost of power generated by the
out-of-service unit. For example, assuming a power cost differential of
10.0 mills/kWh for three different downtime periods of 1, 3, and 6 months
1-5
-------
with a capacity factor of 50 percent, the following additional capital
investments would be required:
Downtime Period Downtime Replacement Power Costs
(Months) ($/kW)
1 4
3 11
6 22
New coal-fired plant cost of power would be approximately 60 mills/kWhr with
half the cost being fixed costs and half the cost being fuel and consumable
costs. For post combustion technologies the downtime replacement power cost
1s less of a factor than for In-situ technologies. Constant dollar
calculations are based on standard return on investment (i.e., annuity)
calculations without consideration of tax incentives (e.g., accelerated
depreciation, investment tax credits) or allowance for funds used during
construction (AFDC). The cost calculations include a state and federal
income tax rate of 38 percent.
The costs presented in the appendices are in current 198B dollars and a
30-year book .life. To approximate the total level1zed busbar cost of power
in constant dollars, divide the current dollar costs by 1.75.
SUMMARY OF RESULTS
Table 1-3 and Figures 1-1 through 1-3 summarize for each technology the
range of cost estimates developed in Table 1-2 using the high and low case
values. The most representative value, the base case, is shown on the
figures for each technology as a mid-way point on the bar graphs. This is
to show the technology sensitivity to variation in boiler and coal
characteristics and that there is no single "winner" for all retrofit
applications.
Only those costs which were developed using the IAPCS cost model were
presented in this section for consistency. Cost estimates for other
technologies which were obtained from other references are presented in the
respective technology sections.
1-6
-------
Sensitivity case cost estimates developed using the IAPCS cost model
are also presented in the appendices. The major cost parameters were
varied for the sensitivity analysis. Major cost parameters differ for the
different technologies. For example, FGD costs are very sensitive to
retrofit factors, coal sulfur content, capacity factor, and boiler size,
while coal switching is mainly a function of fuel price differential and
percent reduction required. A list of sensitivity case parameters for
different technologies is summarized in Table 1-3.
Figures 1-1 through 1-3 present cost estimates for both low and high
cases for capital, levelized annual, and unit costs. Costs as well as
pollutant removal efficiencies vary for different technologies. These two
factors should be balanced in choosing one technology over another and
determining the cheapest technology for meeting acid gas removal
requirements for a given boiler and coal characteristics.
In this study both high and low sulfur coals are switched to a 0.9
percent West Virginia bituminous coal. Therefore, for high sulfur coal (low
case) over 80 percent S02 removal is achieved, while for low sulfur coal
(high case) the removal value is less than 10 percent. Because of a very
low removal efficiency due to switching from one low sulfur coal to another
low sulfur coal with less than 10 percent SO* removal, the unit cost (dollar
per ton of S02 removed) resulted in a very large number (the division
denominator was a very small value for tons of S02 removal). The AFBC and
IGCC costs presented are for new systems. The costs for these two
technologies are much higher than for other technologies presented in this
report because pulverized coal boiler costs (equivalent to AFBC and IGCC)
are not included with the other technologies. For FSI, it 1s assumed that
70 percent S02 removal can be achieved with humidification and that existing
ESPs are adequate in size and can be reused. Therefore the major cost items
are sorbent preparation and modification of the existing furnace for sorbent
injection.
SCR costs are much greater than other NOX removal technologies. This
is mainly due to the initial as well as the replaced catalyst cost.
However, unlike other N0x removal technologies, SCR can achieve more than 80
percent NO removal.
1-7
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TABLE 1-3. SUMMARY OF COST RESULTS - CONSTANT 1988 DOLLARS
Emission Reduction
Technology
Commercial
Fuel Switching and Blending
Lime/Limestone fGD
Lime Spray Drying with
reuse of existing ESP
Lime Spray Drying with
new fabric filter
Lou NO Burners
Overfire Air
Near Commercial
Add-on Controls:
Furnace Sorbent Injection
Natural Gas Reburn
Selective Catalytic Reduction
Advanced Combustion Systems:
Integrated Casi.f ication
Combined Cycle*
Atmospheric Fluldlzed
Bed Combustion*
Capital Costs
Percent
SO
2
2-80
90
76
86
Q
0
70
15
0
95
90
NO
X
0
0
0
0
50
25
0
60
80
60-70
20-50
MU - size in megawatts, XS - coal sulfur content, CF
FPD - fuel price differential.
RF - retrofit
factor.
Lou
20
120
70
UO
B
2
25
10
90
1,710
1,560
-------
3000
2800
Legend
= Base case vclue
ICCC
to
o
o
<
o
2600
2400
2200
2000
1800
1600
1400
1200
1000
800
600
400
200
CS/8
. 0
LSD+ESP
20 30 40 50 60 70
PERCENT NOX OR SO2 REMOVED
L/LSFGO
Figure 1-1. Capital Costs - Constant 1988 Dollars
1-9
-------
700
600
M
O
u
I
200 —
100 —
Legend
Base cose value
\
10
20
I
30
*O 50 60 70
PERCENT NOX OR S02 REMOVED
Figure 1-2. Levelized Annual Cost - Constant 1988 Dollars
1-10
-------
o
CO
o
o
H
Z
10
20
30 40 50 60 70 80
PERCENT NOX OR S02 REMOVED
100
Figure 1-3. Unit Cost - Constant 1988 Dollars
i-ll
-------
References
1 Maibodi, M., A. L. Blackard, and R. J. Page. Integrated Air Pollution
Control System. Draft Report. U. S. Environmental Protection Agency,
Research Triangle Park, North Carolina. February 1990.
2. Emmel, T. E. and M. Maibodi. Retrofit Costs for SO- and NO Control
Options at 200 Coal-Fired Plants. Draft Report. UT S. Environmental
Protection Agency, Research Triangle Park, North Carolina. December 1989.
3. TAG - Technical Asesssment Guide. (Volume 1). EPRI Report P-4463-SR.
Electric Power Research Institute, Palo Alto, California, 1986.
1-12
-------
SECTION 2
COMMERCIAL TECHNOLOGIES
INTRODUCTION
A variety of commercial technologies are available for reducing SOp and
NO emissions from existing power plants. These technologies include both
A
passive (primarily associated with technologies which generally require 1ittle
or no new equipment) and active (based on installation of new equipment)
controls. For purposes of presentation, these controls are divided into four
main groupings and represent the following six technology areas:
• Passive Controls
- Fuel Switching and Blending (Section 2.1)
- Least Emissions Dispatch (Section 2.2)
• Pre-Combustion Controls
- Physical Coal Cleaning (Section 2.3)
• Combustion Controls
- Combustion Modifications (Section 2.4)
t Post-Combustion Controls
- Throwaway Wet Flue Gas Desulfurization (FGD) (Section 2.5)
- By-product Recovery FGD (Section 2.6)
In addition, a discussion of combining these technologies to further reduce
total plant emissions is presented in Section 2.7.
2-1
-------
2.1 F.UEL SWITCHING AND BLENDING
Description
Switch Ing--
Coal switching represents one of several passive (non-hardware) control
methods for emissions reduction. The sulfur content of coals used in the U. S.
ranges from less than 0.5 percent up to roughly 6 percent. Thus, one
alternative for reducing S02 emissions from coal combustion 1s for power plants
to switch from burning high-sulfur coal to lower-sulfur coals.
Most coal-burning facilities are designed to burn coals within a
specified range of heating value, ash content, and other physical and chemical
properties. Largely as a result of empirical experience gained over time,
boiler manufacturers have developed specialized knowledge of how to design a
boiler for a given coal. As shown in Figure 2.1-1, significant variations in
boiler size and configuration exist among boilers designed for coals wiIn-
different fuel characteristics. Switching an existing boiler to a different
coal, an area in which boiler manufacturers have had relatively little
experience, can result in mismatching plant capabilities and may require major
modifications to the plant. Primary areas of concern include coal handling and
pulverization, combustion kinetics, ash deposition on heat transfer surfaces,
particulate collection, and ash handling. Major coal characteristics and
hardware systems which must be considered in evaluating coal switching are
listed in Table 2.1-1.
The impact of these factors on a given boiler is coal and boiler specific.
In general, switching to a coal of the same rank (I.e., conversion from a
high-sulfur bituminous to a low-sulfur bituminous) will cause fewer problems
than conversion to a lower rank coal. However, other coal properties may still
restrict fuel switching in individual boilers. For example, most low-sulfur
coals have high fusion temperature ashes which do not slag at normal furnace
temperatures, effectively precluding their use in cyclone and wet-bottom
boilers.
Power plant systems other than the boiler may also be affected. Of
concern in evaluating fuel switching as an S02 control option is that low-
2-2
-------
Legend
W • Width
D = Depth
h • Nose / Top Burner Row Distance
H • Furnace Height
WxD
1.08WX 1,080 1.18WK1.08D 1.2BWK 1,240
1.29W X 1.28D
EASTERN
MIDWESTERN
BITUMINOUS
SUB-BITUMINOUS
WILCOX SEAM
TEXAS LIGNITE
YEGUA-JACK80N NORTHERN PLAINS
LIGNITE
Figure 2.1-1. Effect of coal rank on furnace sizing (1).
2-3
-------
TABLE 2.1-1. TECHNICAL FACTORS AFFECTING COAL SWITCHING
Change In
Coal Characteristics
Lower heating value
Higher moisture
Higher ash
Lower sulfur
Higher ash fusion
Higher sodium
and iron content
in ash
Higher volatility
Harder grindability
Resulting Potential
Qperatino Problems
Insufficient coal
handling capacity
Unable to achieve design
steam output
Longer/cooler flames
Lower furnace exit
temperature
Higher gas flow
Increased particulates
in flue gas
Increased solid waste
Lower particulate
collection efficiency
Incompatible with cyclone
and wet-bottom furnaces
Increased slagging and
foul ing
Changed heat transfer
characteristics
Heating and potential fires
in coal handling equipment
Insufficient pulverizer
capacity
Possible Solutions
Enlarge coal handling
equipment
Derate capacity
Derate capacity
Increase/modify boiler
heat transfer surface
area
Increase fan capacity
Increase soot blowing
Modify boiler convective
section heat transfer
area
Increase ESP plate area
Modify ash handling and
disposal systems
Increase ESP plate area
Install flue gas
conditioning
Use different coal
Increase soot blowing
Use ash additives
Accept higher forced
outage rates
Change boiler tube
distribution in
furnace and convective
sections
Modify pulverizers,
silos, and other coal
handling equipment
Increase pulverizer
capacity
Derate boiler capacity
2-4
-------
sulfur coals generally have higher ash resistivity. As a result, ESP
performance is likely to be poorer than with high-sulfur coals. This may be
especially critical in boilers equipped with older ESP's which are only
marginally in compliance with existing particulate standards.
Blendlng--
Coal blending to reduce the average sulfur content of coals is another
strategy for compliance with existing SO- emission limits. Depending on
plant-specific requirements, "blending" can denote either the use of multiple
coals by a single plant even though each boiler at the plant may burn a single
coal or the use of multiple coals in a single boiler. In the first case,
several distinct coal piles are maintained for supplying individual boilers. A
central 'coal receiving facility and certain other common equipment may be used
to handle all coals used at a plant, but most of the equipment connecting the
coal pile to the boiler is unit specific. The major limitation on retrofit
application of this form of coal blending is the availability of space for
additional coal piles.
The second form of blending involves the use of two or more coals in a
single boiler. If the individual coals have significantly different combustion
or ash characteristics (e.g., reactivity, ash fusion temperature, or base/acid
ratio), failure to adequately blend the coals into a uniform mixture can result
In rapid fluctuations in fuel composition being fed to the boiler and cause
major plant operational problems. Although a number of power plants receive
coals from different suppliers, in most cases coals purchased by an individual
plant come from the same coal producing area and have similar combustion and
ash characteristics. In these cases, blending operations may be fairly simple
and relatively inexpensive.
When significant differences in fuel quality exist or when continuous
monitoring requirements on plant emissions impose short averaging times,
additional coal-handling equipment (stack-out equipment, storage piles, reclaim
hoppers, conveyors, and belt scales) must be added to the plant to assure
adequate mixing of the individual coals. Capital and operating costs of this
equipment can be significant, as evidenced by Detroit Edison's $350 million
2-5
-------
coal blending facility at its Monroe Station. The facility has the capability
of blending up to 10,000 tons/hr of high-sulfur bituminous and low-sulfur
subbiluminous coals.
Combustion characteristics of coal blends cannot be accurately estimated
as simple linear averages of the individual coals used. This is especially
true for ash slagging, fouling, and resistivity which are influenced by
eutectlcs formed from the mineral matter in the Individual coals in the blend.
The only way to be sure a given coal or coal blend can be satisfactorily burned
in a given boiler is to conduct a test burn of several days to several weeks in
the actual boiler. Preliminary testing of candidate blends in a pilot-scale
combustor can be useful in better defining potential problems prior to full-
scale tests.
Applicability
Low-sulfur coals (<1 percent sulfur) are available in several regions of
the U. S.: central and southern Appalachia, Powder River Basin, southern
Wyoming, and Colorado/Utah (Figure 2.1-2). Low-sulfur coals from southern
Appalachia are expensive to mine and limited In quantity, and are therefore
unlikely to be used as an alternative to high-sulfur coal in existing boilers.
Basic coal characteristics for the remaining low-sulfur coals and for high-
sulfur coals from northern Appalachia and the Illinois Basin are presented in
Table 2.1-2. On a pounds of SO. per million Btu basis, low-sulfur coals
generally have one-quarter to one-half the sulfur content of the high-sulfur
coals.
Central Appalachia coals are used widely in the eastern U. S., accounting
for virtually all of the coal used 1n Virginia, North Carolina, South Carolina,
southern West Virginia, eastern Kentucky, and Tennessee. Significant amounts
of central Appalachian coal are also used in Ohio and Michigan to meet SIP
requirements for low-sulfur coal. Estimated reserves of low-sulfur coal in the
region are 23 billion tons. Production in 1982 was 118 million tons. Coal
seams In the region are generally thin (2.5 to 7 feet thick). Mining and
reclamation in the region is difficult due to the area's mountainous topography
and result in high mining costs. Coal prices FOB mine over the last few years
have ranged from $30-40/ton (2,3).
2-6
-------
ro
i
1 Northern Appalachia
2 Central Applachia
3 Southern Applachia
4 Illinois Basin
5 Western Interior
6 Gulf Lignite
7 Plains Lignite
8 Powder River Basin
9 Southern Wyoming
10 Colorado/Utah
11 Arizona/New Mexico
12 Washington
SOURCE Bwgy Ventures Analysis. Inc., COALCAST Quartorty Report Service, 1983.
Figure 2.1-2. Coal supply regions (1).
-------
TABLE 2.1.Z TYPICAL DELIVERED COAL PROPERTIES FOR SELECTED SUPPLY REGIONS"
Property
Rank
Higher Heating range
Value, Btu/lb. average
Sulfur, X range
average
Moisture, X
Ash, X range
average
Crindabilltyb
Ash Fusion Temperature,
Fluidity, -F
Northern
Appalachia
Bituminous
11,000-13,000
12,138
0.7-5.0
2.5
5-10
7-15
12.8
50-100
2.100-2,800
Central
Appalachia
Bituminous
11,500-13,500
12,252
0.6-3.5
1.1
5-10
6-15
10.9
40-90
2,10Q-2,800
Illinois
Basin
Bituminous
10,000-12,000
11,114
1.0-3.5
2.8
5-15
8-15
10.5
50-70
1.900-2,400
Powder River
Basin
Subbi tumi nous
8,200-9.700
N/A
0.3-0.7
N/A
20-30
5-9
N/A
40-55
2,100-2,300
Southern
Wyoming
Subbi luminous
9,000-10,500
N/A
0.4-1.0
N/A
15-20
7-11
N/A
40-55
2,100-2,400
Colorado/
Utah .
Bituminous
10,000-12,500
11,130
0.3-0.8
0.5
7-20
6-13
9.3
45-50
2,200-2,600
*See Reference 1.
bHardgrove grindability index (AST* 0-409).
-------
Western low-sulfur coals, especially those from the Powder River Basin,
are currently being shipped to utilities in the western, south-central, and
midwestern U. S. Estimated reserves of low-sulfur coal exceed 140 billion
tons. Total production in 1982 was 157 million tons. Coal seams range up to
over 100 feet thick in the Powder River Basin where mining is by large pit-type
surface mines. Seams in southern Wyoming, Colorado, and Utah are generally
thinner and are recovered by both surface and underground mines. Coal prices
FOB mine have ranged from less than $6/ton in the Powder River Basin up to
$30/ton in Utah (1,3).
Two major questions exist regarding the future price of low-sulfur coals:
the impact of acid rain policies on 1) minemouth low-sulfur coal prices
(especially in central Appalachia) and 2) transportation rates from the mine to
the consumer. Several studies based on computer modeling have estimated that a
major acid rain control program could increase the demand for central
Appalachia low-sulfur coal by up to 50 million tons over otherwise projected
levels and increase minemouth prices by as much as $15 per ton (4). Minemouth
costs of low-sulfur western coals are not expected to increase as much due to
region's current excess mining capacity and larger reserves. However, regard-
less of coal prices FOB mine, railroad transportation rates may increase the
delivered cost of low-sulfur coal and decrease the attractiveness of fuel
switching. Other analysts argue that recent structural and technological
changes in the coal and railroad industries are not captured in these modeling
efforts, and that price increases will be less than projected (1). For
example, increases in labor productivity within the coal industry and the
declining demand for metallurgical coal by the steel industry have both worked
to keep the cost of coal from central Appalachia below forecast prices.
Performance
In 1985 the 31 states adjoining or east of the Mississippi River contained
230,000 MW of coal-fired capacity. Of this total, roughly 171,000 MW had S02
emissions in excess of 1.2 Ibs/million Btu. Total annual SO^ emissions from
these units were 12.6 million tons. An estimated 85,000 MW of this total are
believed to be technically capable of converting to low-sulfur coal. The
remaining 86,000 MW include cyclone and wet-bottom furnaces unable to use low-
2-9
-------
sulfur coals due to high ash-fusion temperatures and minemouth plants with
limited transportation alternatives. If all of these units are converted to
lower sulfur coals, annual SO- emissions would decrease by roughly 4-7 million
tons depending on the sulfur contents of the coal fired in the units prior to
being converted and utility dispatch decisions. The resulting increase in the
demand for lower sulfur coals would be roughly 125 million tons per year.
The actual extent of fuel switching will partially depend on the nature of
the SO* reduction program mandated. Key factors impacting fuel switching
benefits at a given plant include the delivered price of low-sulfur coal, the
sulfur and Btu content of the new fuel, the relative capital cost of FGD
retrofit versus plant modifications for fuel switching, plant-specific
considerations such as plant layout and projected capacity factors, and the
expected ease of implementing a major SO. control program throughout an
integrated utility system. Potential SO. reductions at a typical plant range
from 50-80 percent.
Cost
Fuel switching and blending are generally assumed to have low capital
cost. While in many instances the capital cost for conversion between two
coals may be small, this assumption is not true in every instance.
Unfortunately, many analyses of SO- retrofit costs have Ignored the technical
limitations and capital costs associated with coal switching.
The technical and economic feasibility of switching fuels in a given
boiler are specific to the boiler design and location. Where switching is
technically feasible, economics, will depend on the cost of coal delivered to
the plant, the capital and O&H costs associated with necessary plant
modifications, the cost of replacement power associated with plant derating
when a different coal 1s used, and the comparative sulfur contents of the
existing and candidate coals. In this study using the IAPCS cost model (5),
estimated capital costs are for particulate control improvements and fuel price
differential (FPD). Particulate control impacts were evaluated by considering
the altered particulate loading and particulate resistivity associated with the
replacement coal as compared to original conditions. ESP upgrades were costed
2-10
-------
by calculating additional plate area in the presence of SCL gas conditioning to
maintain current PM emission rates. The premiums of $5 and $15 per ton were
used to span the range of fuel cost increased associated with low sulfur coal
demand. Using the EPRI cost procedures (6), capital costs include
preproduction costs estimate as 25 percent of full capacity fuel cost for one
month. This results in a higher capital costs for $15 FPD versus the $5 FPD.
The IAPCS costs do not include the cost impact of additional coal
receiving/storage/handling facilities (if needed) and the cost impact of boiler
derate due to pulverizer capacity and boiler fouling, slagging, and erosion.
The majority of the capital costs are due to coal inventory costs
(preproduction costs).
The following table provides the range of values for capital cost,
annualized cost, and cost per ton of pollutant removed estimated using the
IAPCS cost model. The lower capital and annualized costs are for a large unit
with high capacity factor and $5 FPD, while the higher capital and annualized
costs are for a small unit with low capacity factor and $15 FPD. Figures 2.1-3
through 2.1-5 show these costs as a function of boiler size and fuel price
differential. Appendix A contains tables for estimating coal switching costs
as a function of size, coal sulfur content, capacity factor and FPD.
Range
Capital Cost ($/kW) 19.1 to 43.1
Annualized Cost (mills/kWh) 5.4 to 17.2
Cost Per ton of SO, Removed ($/ton) 213.7 to 20,354.1
e.
The delivered cost of low-sulfur coals used in these calculations may not
include sufficient fuel premiums to account for rapid increases in low-sulfur
coal demand as a result of "acid rain" legislation.
Table 2.1-3 summarizes key factors influencing the decision to switch coal
versus retrofit a wet scrubbing system or some other form of active emission
control technology.
2-11
-------
Legend
fuel Price Differentiol = 5 $/ton
"uel Price Differential = 15 $/ton
Sulfur Content = 1 or 2 %
Cooocity Factor = 50 %
o
a
18
100
300
500
MW
700
900
i.lOO
1.300
Figure 2.1-3. Coal Switching - Capital Cost, Current 1988 Dollars
oc
111
Q.
V)
I
1 D -
15 -
14 -
13 -
12 -
1 1 -
10 -
9 -
8 -
7 -I
6 -
5 -
10
Legend
• Fuel Price Differentiol = 5 J/ton
+ Fuel Price Differentials 15 S/toi
Sulfur Content = 1 or 2 7,
Capacity Factor = 50 %
^^^^
1
0 300 500 700 900 1.100 1.300
MW
Figure 2.1-4. Coal Switching - Levelized Annual Cost,
Current 1988 Dollars
2-12
-------
0.0
is -
18 -
16 -
14 -
13 -
12 -
1 1 -
10 -
9 -
8 J
7 -
6 -
5 -
4 -
3 -
1 -
T
~"~~ --—^^ .
i^^^
Legend
Stifir Content . 1%
• Fuel Price Differential = 5 $/ton
+ Fuel Price Differential = 15 $/ton
Stitur Content - 2%
o Fuel Price Differential = 5 $/ton
A Fuel Price Differential = 15 $/ton
Capacity Factor = 50 %
i i i i
1
100 300 500 700 1000 1300
MW
Figure 2.1-5. Coal Switching - Cost Per Ton of, SO2
Removed, Current 1988 Dollars
2-13
-------
TABLE 2.1-3. FACTORS INFLUENCING THE DECISION TO SWITCH
OR SCRUB TO REDUCE SULFUR DIOXIDE EMISSIONS
Factor
Coal
Switching
Retrofit
Scrubbers
High-sulfur coal cost
Boiler type
High-sulfur coal supply
Low-sulfur coal transportation costs
Powerplant age
Difficulty of scrubber construction
Existing coal delivery facilities
Located In high-sulfur coal
high*
dry bottom
not restrictive
low
old
high
rail, barge
low*
cyclone
restrictive
high
new
low
truck, conveyor
producing state
Required SO. emission reduction
Capacity margin
System costs of derating
Load growth outlook
Utility financial condition
no
low
high
low
low
weak
yes
high
low
high
high
strong
Source: Reference 1.
*This table 1s read in the following manner: If the current cost of high
sulfur coal delivered to the plant is high, the utility is more likely to
choose coal switching. If the cost of the high sulfur coal is low, the
utility is more likely to choose a scrubber retrofit.
2-14
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2.2 LEAST EMISSIONS DISPATCH
Description
One of the key operational decisions facing a utility (or group of
interconnected utilities) on day-to-day basis is which plants to operate or
"dispatch" to satisfy the demand for electricity at any given time. Presently,
utilities dispatch plants in accordance with state regulatory guidelines to
minimize the cost of generation, subject to physical constraints such as
individual plant availability, flexibility in adjusting to variations in
electrical demand, and plant maintenance requirements. If environmental
regulations were changed to limit system-wide emissions of SO- or if emissions
were taxed to reflect the "cost" of emissions, the economics of plant dispatch
would be changed and a different dispatch pattern would potentially occur.
This alternate dispatch pattern has been referred to in various publications as
"least emissions dispatch" (LED) and "environmental dispatch".
Several different LED concepts have been described in the literature. Key
differences in these concepts are the time period (temporary vs. permanent) and
geographic area affected. One proposal is to use LED to reduce annual
emissions over a large area. In this approach, maximum use is made of nuclear
and hydroelectric capacity and low-emitting fossil-fuel plants (natural gas,
low-sulfur oil and coal, and plants with scrubbers for SCL reduction) in
preference to plants burning hi.gh-sulfur fuels. Interregional transfers of
power are also used, including the importation of electricity from Canada.
The other major LED approach would reduce SO- emissions only during
periods of adverse air quality in environmentally sensitive areas. For
example, LED could be used during periods of air stagnation in the Adirondacks.
The objective of this approach is to reduce ambient concentrations of S02 in
selected areas as opposed to reducing total emissions of all plants in a large
region. To be effective, this approach requires knowledge of source/receptor
relationships affecting pollutant transport and transformation. The only known
use of episodic LED has been in the Los Angeles air basin where power plant
operations were controlled to reduce NOX emissions during periods of poor air
quality. This effort is reported to have been a failure and was discontinued
(7).
2-15
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Applicability and Performance
Because every utility and region of the U.S. is different, assessment of
the potential and cost effectiveness of LED for reducing SO- is utility
specific. For utilities operating both large, thermally efficient plants
equipped with FGD systems and smaller, less efficient plants without pollution
control systems, least cost and least emissions dispatching may result in
operation of the same plants. For utility systems which can import power from
other geographic regions or which depend on plants that are not equipped with
pollution control equipment for base-load power generation, LED may signifi-
cantly alter dispatch decisions. Also, the utility's capacity factor will
influence whether LED can be effective. Utilities with low capacity factors
may be able to alter dispatching significantly and thus appreciably reduce SO-
emissions. Utilities with high capacity factors have limited opportunity
to shift generation between plants. To the extent low-sulfur coal plants are
dispatched In preference to high-sulfur coal plants, the impact of LED in coal
markets will be similar to coal switching.
The greatest potential for LED is in a large utility or interconnected
utility pool having multiple plants, using a variety of fuels, and having a
relatively low system-wide capacity factor. Although the concept of LED has
been presented in a number of studies, very little analytical work has been
done to quantify its cost effectiveness or the magnitude of potential emission
reductions. Given the requirements of state regulatory agencies to minimize
the cost of electricity, such analyses will be required before LED can be
incorporated into or replace the existing dispatch process.
Cost
LED does not involve capital expenditures and only the differentials in
fuel and variable O&M costs need to be considered in estimating the cost of
LED. For example, the cost differential
per kilowatt-hour from dispatching a plant burning $41/ton central Appalachia
coal with 0.7% sulfur versus a $35/ton northern Appalachia coal with 2.4%
sulfur (assuming a 10,000 Btu/kWh heating value and equal O&M costs for both
Z-16
-------
plants) Is $0.0025 ($6 coal price differential divided by 24 million Btu per
ton). This equates to a 30-year current 1988 dollar cost per ton of SO-
removed of $334.
Estimating the cost-effectiveness of LED for an entire utility network is
analytically similar to "bubbling" regional emissions. The major analytical
difference with LED is that the focus is on altered dispatch to minimize SO-
rather than to minimize cost. As with fuel switching, the cost of the first
ton of SO. reduction using LED is small. As the number of altered dispatches
Increases, the incremental cost of LED can increase significantly.
Two studies have been published on the cost-effectiveness of LED on a
regional basis. These studies were conducted by the Northeast States for
Coordinated Air Use Management (NESCAUM) (7) and the Wisconsin Public Service
Commission (8) as part of EPA's State Acid Rain (STAR) program.
The NESCAUM study examined SO- and NO control alternatives for the region
containing New York, New Jersey, and the six New England states. This
analysis, based on a regionalized version of EPA's Advanced Utility Simulation
Model (AUSM), concluded that the region's utilities could reduce SO- emissions
by 35 percent (158,000 tons) and NO emissions by 41 percent (51,900 tons)
^
through use of LED at an average current 1988 dollar cost of $4,893/ton.
The Wisconsin study estimated that If 1) all power plants in the state
were centrally dispatched and 2) significant conservation efforts are
implemented (resulting in relatively low system capacity factors), LED could
reduce SO- emissions by roughly 15 percent (73,400 tons) at an average current
dollar cost of $494/ton of S02- At higher system capacity factors, the SO-
reduction potential of LED is reduced.
The large difference in projected cost from these two studies indicates
the regional- and utility-specific differences in LED effectiveness, as well as
the significant difference in cost associated with various emission reduction
targets (15 percent in Wisconsin versus 35 percent in New England).
2-17
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2.3 PHYSICAL COAL CLEANING
Description
Run-of-mine coal consists of Individual particles containing varying
amounts of organic and Inorganic matter. The Inorganic materials are largely
sedimentary rock (e.g., shale, claystone, sandstone, and limestone) and pyritic
sulfur. Physical coal cleaning separates the combustible organic material
(i.e., coal) from the non-combustible impurities. Although the primary
economic Impetus for most existing coal cleaning Is removal of ash, significant
benefits also accrue from the removal of sulfur. Of the total coal consumed by
power plants, about 42 percent 1s cleaned using coal cleaning processes (9).
For most medium- and high-sulfur bituminous coals, pyritic sulfur which
can potentially be removed by physical separation methods accounts for 30 to 70
percent of the total sulfur. Most of the remaining sulfur is associated with
the organic structure of the coal and is not separable by physical means. The
upper limit on sulfur removal depends on the physical liberation of pyrite from
the organic structure during coal crushing prior to cleaning. Sulfur contents
of some Northern Appalachla coals (e.g., the Freeport and Kittaning seams of
Pennsylvania) can be reduced by 50-60 percent using currently available
cleaning methods, but removals of 30 percent or less are common for most coals.
Most commercial techniques rely on differences In specific gravity to
accomplish separation.(10) For example, clean bituminous coal has a specific
gravity (s.g.) of 1.2-1.3; shale, 2.0-2.7; and pyrite, 4.8-5.0. Selective
adhesion, magnetism, and surface tension are also used.
In cleaning processes based on specific gravity differences, the raw coal
is separated into float and sink fractions by submerging it in a fluid with a
specific gravity between that of the coal and the inorganics. In some
technologies, such as jigging and tabling, the differences in specific gravity
are combined with mechanical separation techniques to enhance the recovery of a
clean coal fraction. However, the clean coal float will always contain some
impurities, and the rejected sink fraction will contain some coal. There are
two reasons for this. First, inherent inefficiencies in separating processes
cause a fraction of heavy particles to be misplaced in the float instead of the
2-18
-------
sink, and vice versa for some light particles. Second, the organic and
inorganic matter cannot be totally separated because they frequently occur in
the same particles.
Yields and qualities of clean coal achievable for different specific
gravities and particle top sizes can be estimated from laboratory washability
data. Figure 2.3-1 shows washability data for two coals. The coal shown on
the left is easily cleaned because, as shown on Curve A, the ash content of
incremental amounts of coal recovered increases rapidly as the specific gravity
of separation exceeds a given level, in this case roughly 1.5 (Point X). As
shown on Curve B, the cumulative ash content of the recovered coal at this
specific gravity is less than 10 percent versus an original ash content of
roughly 30 percent (as plotted on Curve B at 100 percent recovery). The
cumulative ash content of the sink material (Curve C) at s.g.=1.5 is roughly 65
percent. The right side of Figure 2.3-1 represents a difficult-to-clean coal
for which there is no easily definable point of separation. Commercial
cleaning equipment is most efficient when separations are easily made.
Because particles of different sizes behave differently when suspended in
a fluid, technologies have been developed to clean each size fraction
separately. These technologies are based on three general size classifi-
cations:
Fraction Classification Typical Size Range
Coarse greater than 3/8 inch
Intermediate 28 mesh (0.6 mm) to 3/8 inch
Fine 0 to 28 mesh (0.6 mm)
Jigs and dense media (DM) vessels are the most common equipment used for
coarse coal cleaning. A jig feeds raw coal through a column of pulsating
water. These pulsations suspend the lighter coal particles at the top of the
jig while allowing the dense ash and pyrite particles to sink. Dense media
vessels use a suspension of fine magnetite particles (Fe304) and water to float
all material below a certain s.g. After separation, the refuse and clean coal
are screened and rinsed with water to remove the magnetite. Magnetite in the
rinse water is reconcentrated using a magnetic separator. Figure 2.3.1
Coarse coal cleaning accounts for one-half to two-thirds of all coal cleaned in
the U. S.
2-19
-------
rv>
o
A- Easy-to-Clean Coal
WEIGHT RECOVERY OF CLEAN COAL (%)
SPECIFIC
GRAVITY
B-Difficult-to-Clean Coal
WEIGHT RECOVERY OF CLEAN COAL (%)
SPECIFIC
GRAVITY
0 10 2O 30 40 50 60 70 80 90 100
ASH CONTENT (%)
1O 2O 3O 40 SO 60 70 80 90 10O
ASH CONTENT (%)
Curve A
Incremental Ash Content
In the Float
Curve B Curve C
Cumulative Aeh In the Float Cumulative Ash In the Sink
Near Gravity Material
Figure 2.3-1. Coal washability curves. Reference 11.
-------
Intermediate-sized particles are generally cleaned using concentrating
tables and dense media cyclones. A concentrating table is an inclined,
vibrating surface, equipped with diagonal riffles. Raw coal slurry is fed to
an elevated corner of the table. The riffles act as miniature jigs to stratify
the raw coal. Lighter clean coal is carried by water over the riffles and off
the lower end of the table. Impurities sink between the riffles and are
conveyed to the side of the table by the vibration. DM cyclones use
centrifugal force to separate clean coal from refuse, much as gravity is used
in dense media vessels. Because small particles settle slowly under gravity,
the much stronger centrifugal forces in a cyclone permit larger tonnages of
fine particles to be handled than in a DM vessel. Intermediate coal cleaning
accounts for one-quarter to one-third of U.S. cleaned coal.
Fine coal (less than 28 mesh) is difficult to handle because of problems
1n separating fine particles from the cleaning fluid. The principal commercial
methods for cleaning fine coal are froth flotation and fine coal cyclones,
accounting for 5-10 percent of total production. Froth flotation takes
advantage of the hydrophoblc (i.e., water repelling) tendencies of most coals,
versus ash which is hydrophilic (i.e., water adhering). When air is bubbled up
through a flotation cell, the small coal particles attach to the air bubbles
and rise to the surface.
Coal cleaning plants are frequently classified based on the following
definitions (12):
Level 1 - Crushing and particle sizing only. Little or no cleaning
takes place.
Level 2 - Coarse (+3/8 inch) coal cleaning only.
Level 3 - Coarse (+3/8 inch) and intermediate (3/8 inch to 28 mesh)
coal cleaning. Material less than 28 mesh is not cleaned
and is either incorporated with the product or discarded
depending on its quality.
Level 4 - All size fractions are cleaned, but only a single product
consisting of a blend of cleaned fractions is sold.
Level 5 - This level involves the most rigorous coal cleaning, and
produces two or more usable coal products of different ash
and sulfur contents from one raw coal.
2-21
-------
This numbering Is generally compatible with existing coal cleaning
literature. A schematic flow diagram for a Level 4 cleaning plant 1s shown in
Figure 2.3-2. The quantity of material processed as coarse, intermediate,
and fine coal will depend upon the characteristics of the coal and the desired
product quality. The distribution of Level 2 through 4 cleaning plants by
region and level is shown on Table 2.3-1. Level 1 plants are omitted because
of their minimal effect on ash and sulfur levels of coal. The only Level 5
cleaning plant in the U. S. is the Homer City cleaning plant In Pennsylvania.
Applicability and Performance
Most coal cleaning plants in the U. S. were installed by coal companies to
meet specific ash and sulfur specifications. Approximately two-thirds of the
coal in the eastern U. S. is physically cleaned (12). For the nation as a
whole in 1980, 241 million out of 687 million tons of steam coal were cleaned,
or 35 percent.
Table 2.3-2 provides average sulfur dioxide emission rates for major seams
in the two major high-sulfur coal regions in the U. S. Although not shown on
the table, current coal cleaning plants (designed primarily to remove mineral
matter) generally remove 20-30 percent of the sulfur 1n Illinois Basin coal and
from 10-40 percent of the sulfur from Northern Appalachia coal. This level of
sulfur removal is accomplished while recovering most of the heat content of the
raw coal (>95 percent), thus minimizing the dollar value of coal lost during
cleaning. Estimated emissions from state-of-the-art removal are assumed to be
representative of cleaning in a Level 4 preparation plant. In general,
state-of-the-art coal cleaning is most effective for bituminous coals from
North Appalachia (especially those seams in the Allegheny Formation such as the
Freeport and K1ttan1ng seams) where sulfur reductions near 50 percent are
possible and the Illinois Basin which contain significant amounts of
macroscopic pyrite. Additional sulfur removal is possible, but a substantial
portion of the heating value is lost in the reject material.
2-22
-------
ro
i
u>
RAW
FEED
RRFAKPR
1
WET SCREEN AT
- 3/8 INCH
*
WET SCREEN AT
- 28 MESH
j
FROTH
FLOTATION
*
Raw Coarse
Coal
Intermediate Raw
Coal
CteanRne
Coal
fc| p 1 C
^L
Y
POND
JIG OR HEAVY
MEDIA WASHER
DENSE MEDIA CYCLONES
OR TABLES
VACUUM FILTER
RPPIJCP CM TPR
Clean Coarse MECHANICAL
Coal DEWATERING
Clean Intermedate MECHANICAL
Coal DEWATERING
^ 1 H ^^^^M»
S^\ THERMAL
^\J DRYER ~~
^
©CLEA
«b 4 * L. 4L. •!_ 1 »
Determine wrtdtner tne plant CO A
has open or closed water circuit .
O = Open
C a Closed
Figure 2.3-2. Level 4 Coal Preparation Plant
Source: Reference 1.
-------
TABLE 2.3-1. BREAKDOWN OF COAL CLEANING PLANTS
BY REGION AND LEVEL, I97B
Region
Northern Appalachia
Central Appalachia
Southern Appalachia
Midwest
Other
Total
Percent of Total
Total Plants3
128
239
25
53
19
464a
Level 2
35
47
6
23
--
.
Ill
25
Level 3
52
71
10
22
--
k
155°
35
Level 4
41
121
9
8
--
K
179°
40
Source: 1979 Keystone Coal Industry Manual. "Directory of Mechanical Coal
Cleaning Plants."
uDoes not include 12 pneumatic preparation facilities.
Only Includes plants in the four major regions.
TABLE 2.3-2. POTENTIAL REDUCTIONS IN SULFUR DIOXIDE EMISSION RATES FROM
PHYSICAL COAL CLEANING
Northern Appalachla
Pittsburgh
Sewickley
Freeport
Kittaning
Illinois Basin
Illinois 6/Kentucky 11
Illinois 5/Kentucky 9
i
Raw
Coal
5.7
6.2
3.9
4.8
6.5
6.5
State-of-
the-Art1
4.2
4.8
2.2
2.6
4.1
4.4
Maximum .
Achievable
3.3
3.8
1.3
1.7
3.6
3.9
-Based on 90% Btu recovery with 1 1/2" x
Based on 14 mesh x 0 feed at s.g. = 1.4.
x 0 feed.
Teed at s.g. = 1
Source: Reference 14.
2-24
-------
Cost
Physical coal cleaning has many economic benefits besides sulfur removal,
although they can be difficult to quantify (15). These include reduction in
coal transportation costs; reduced coal handling, pulverization, and storage
requirements; generally improved boiler performance; smaller ash handling and
waste disposal volumes; and reduced costs for post-combustion flue gas
desulfurization (FGO) to meet state and federal emission limits. For utilities
having trouble meeting particulate emission limits, this lower ash content can
bring them into compliance and avoid expensive upgrading of ESPs. While these
savings can be significant, they are hard to document quantitatively and vary
from plant to plant. "Typical" cost benefits from coal cleaning have been
estimated at $2.82/ton, but can range from $0.54 to over $9/ton (10). Cost
savings associated with use of coal cleaning to reduce FGD requirements are
discussed in Chapter 5. Other benefits, such as improved boiler performance
and reduced coal handling costs, are very boiler- and site-specific and are not
included in the costs discussed below.
The three major components of coal cleaning costs are the initial capital
equipment, operation and maintenance, and the lost heating value of discarded
coal. The most important variables in determining these costs are the
characteristics of the raw coal and the desired level of cleaning (which in
turn determines the selection of cleaning technology and the quantity of Btu
rejected).
Of the total capital costs, roughly one-third are associated with the
cleaning equipment per se. These costs (including equipment delivery and
installation) range from around $2,200 per ton of raw coal per hour for a jig
or DM vessel processing coarse coal to nearly $22,000 per ton per hour for a
froth flotation cell for cleaning fine coal. Another third is for structural
steel, electrical, engineering and construction fees, and other costs that are
generally proportional to the size and complexity of the project design. The
remaining third covers items such as refuse ponds, site preparation, conveyors,
silos, and load out facilities for which costs are relatively independent of
the cleaning level. On a total plant basis, total capital cost estimates range
from around $11,000 per hourly ton of input capacity for a Level 2 plant to
2-25
-------
near $43,000 per hourly ton for a Level 4 plant Incorporating froth flotation
for fine coal cleaning.
Operating costs increase substantially with higher levels of cleaning, as
well as with coal washability. The operating cost for a coal circuit designed
for cleaning intermediate coal (3/8" x 28 mesh) may be double on a per ton
basis the cost of a circuit for coarse coal (+ 3/8"). More complex plants
require more personnel and greater training in addition to power, supplies and
maintenance. Magnetite costs can be very important in dense media plants. The
heating value of the coal that is thrown away as refuse can be the highest cost
of coal cleaning, and 1s both a function of the level of cleaning and the coal
washability.
As a result, the cost of sulfur removal using coal cleaning varies widely
depending on raw coal quality and user specifications. Table 2.3-3 shows
ranges of capital costs, O&M cost, and Btu loss for several Northern
Appalachian and Illinois Basin coals on a life-cycle basis (i.e., per ton of
cleaned coal). Busbar costs and cost effectiveness are also shown. These
values assume Level 2 and Level 4 cleaning of a raw coal with a sulfur content
of 3.5 percent. Btu recoveries were 90-96.5 percent, weight recoveries 73-90
percent, sulfur reductions 17-62 percent, and raw coal cost $1.30/million Btu.
Although Level 2 cleaning has lower capital and annualized costs than Level 4,
costs per ton of S02 removed Level 4 are frequently lower because of higher
sulfur removal rates. Cost per ton of S02 removed for lower sulfur coals
generally increase due to the decreased amount of removable pyrite in the raw
coal.
2-26
-------
TABLE 2.3-3. ECONOMICS OF CONVENTIONAL PHYSICAL COAL CLEANING
$/Ton of Clean Coal Current 1989 Dollars
Capital 0.54 - 2.43
0 & M 3.79 - 7.84
Btu Loss 1.50 - 6.50
Total 5.83 - 16.77
Busbar Cost (mills/kwh) 2.7 - 6.2
$/ton of S02 325 - 595
2-27
-------
2.4 NOX COMBUSTION CONTROL TECHNOLOGY
The production of NO during combustion of coal (or other fuel) is
dependent on boiler design, size, and operating parameters as well as fuel
composition. By modifying flame geometry and the rate of air/fuel mixing,
thermal oxidation of atmospheric and fuel nitrogen can be reduced.
Specifically, "staging" of combustion to provide a fuel-rich (i.e., reducing)
zone during the initial, high-temperature stage of combustion will favor
formation of molecular N- instead of NO. Additional (i.e., secondary) air is
added after the initial high-temperature reaction to complete combustion of
unburned fuel. Secondary combustion air is added at a point where much of the
fuel-bound nitrogen has been reduced to N2 and flame temperatures are too low
for thermal NOX formation.
Wall-fired wet-bottom and cyclone units are large generators of NO due to
their intense combustion conditions. Tangentially-fired units with their
longer flames and cooler combustion temperature produce less NOX- Because of
the wide variations in boiler characteristics and NO levels, the applicabi-
lity of combustion modification technologies for reducing NO is boiler
specific. This section discusses the two major combustion modification
technologies for NO control: low NO burners and overfire air.
Description
Low NO Burners--
Low NOV burners are designed to reduce NO production (typically from wall
A A
fired boilers) by controlling air/fuel mixing. This changes combustion
reactions in several ways. First, substoichiometric amounts of air'are used
during initial combustion. This results in fuel nitrogen forming molecular
nitrogen (N-) rather than NO. Second, due to controlled addition and mixing
of secondary air, local zones of excess air that promote N0x formation are
minimized. Third, Increasing combustion times and allowing for more heat
transfer from the flame results in a cooler flame and reduced production of
thermal NOX.
2-28
-------
The controlled air/fuel mixing is achieved by using separate air registers
within the burner. Part of the air enters with the fuel and another part
enters through an annulus surrounding the central fuel/air stream. The
remaining air is injected through a second, outer annulus. Separating the air
and fuel delays fuel/air mixing, thus "staging" the combustion process and
creating the low NO conditions described above.
The same effect can be achieved on tangent1ally-f1red boilers, but a
different, somewhat more expensive burner design 1s required. However, the
diagonal firing pattern and larger furnace dimensions of most tangentially-
fired boilers result in lower flame temperatures and less NO formation than
A
wall-fired boilers, thus the potential NO reduction by low NO burners is
A - A •
reduced. Because of this, overfire air is the predominant commercial NO
control technique for tangentially-fired boilers. Therefore, the subsequent
low NO burner discussion will focus on wall-fired units only.
A
Overfire Air--
Another fundamental method of controlling flame stoichiometry, in addition
to (or in place of) the localized combustion staging of low NO burners, is
A
combustion staging within the furnace volume using overfire air (OFA) ports
With this approach, 15 to 20 percent of the required combustion air is diverted
from active burners to OFA ports located above the top row of burners. The
active burners operate fuel-rich, providing a reduction in both thermal and
fuel NO generation. The unburned fuel escaping the fuel-rich flame zone burns
higher in the furnace where the diverted combustion air is mixed. An overfire
air system includes the OFA ports (involving penetrations in the furnace wall),
additional duct work, potentially a separate fan, and dampers/air flow
controls.
Applicability
The applicability and effectiveness of low NOX burners and OFA, especially
for retrofit cases, are dependent on a number of site-specific parameters, some
of which are listed in Table 2.4-1. For example, NOX emissions increase with
increasing coal nitrogen and oxygen contents.
2-29
-------
TABLE 2.4-1. SITE-SPECIFIC PARAMETERS AFFECTING NOV EMISSIONS FROM
COMBUSTION MODIFICATION CONTROLS
Furnace Characteristics - Manufacturer
- Size (width, depth, height)
- Heat Input (Btu/hour) ,
- Heat Release Rate (Btu/hour-ft')
- Firing Type (tangential, wall, cyclone)
- Bottom Ash Removal (slag, dry)
- Number, Spacing, and Size of Burners
- Nose/Top Burner Row Distance
- Gas Velocity
- Furnace Exit Temperature/Attemperation
Margins
- Windbox Size
- Excess Fan Capacity
Coal Characteristics - Heating Value
- Composition (N, 0, H.O, S, Ash, wt. %)
- Ash Fusion Temperature
- Ash Slagging/Fouling Indices
- Volatility
- Variability
Furnace design, size, and dimensions also affect NOX emissions, with low heat
release rate furnaces generally exhibiting lower NOX emissions. Specific
reviews of the applicability of low NO burners and overfire air are presented
below.
Low NOX Burners--
Since the 1971 NSPS, low NO burners have been installed on more than
40,000 HW of boiler capacity. Most of these boilers have been new wall-fired
units with circular burners. An additional 30,000 MW of pre-NSPS wall-fired
boilers are believed technically capable of being retrofit with low NOX
burners. This 30,000 MW represents approximately 20 percent of nation-wide N0>
emissions from pre-NSPS utility boilers (16).
Retrofit of low NO burners to existing boilers is currently being
examined 1n several full-scale utility boilers. Low NOX burners are generally
2-30
-------
larger than conventional burners and require more precise control of fuel/air
distribution. Their performance depends partially on increasing the size of
the combustion zone to accommodate longer flames while avoiding interaction
with flames from other burners. Because of this, low NO burners are expected
to be less effective when retrofit on relatively small furnaces.
Low NOX burners are not currently available for dry-bottom cell-fired and
roof-fired boilers, cyclone boilers, and wet-bottom boilers. Cell burners
consist of two or three burners in a single cluster and were designed
to produce an intense, high-temperature flame. Babcock & Wilcox is currently
conducting a low NO burner development program for cell burner applications
for the Electric Power Research Institute (17). Because of difficulties in
applying low NO burners to cyclone boilers, primary emphasis for reducing NO
^ A
emissions from these units is on reburning technology (see Section 3.7).
In order to retrofit low NO burners, the existing burners must be
replaced. In some cases, some of the water-wall tubes may have to be bent in
order to install the larger low NO burners. Also, low NO burners will
A- A
generally produce a somewhat longer flame; thus, flame impingement on furnace
walls and superheater tubes can be a problem for boilers with high heat release
rates (typical of boilers built in the 1960's) (1). If flame impingement is a
problem, potential solutions include adjusting burner tilt, adjusting
coal/primary air versus secondary air velocities, biased firing of burner rows,
and relocating some superheater tubes. Boilers with very small furnaces may
have to be derated in order to prevent flame impingement at full load.
Other factors to consider in retrofitting low N0¥ burners include fuel
A
characteristics, furnace exit temperature, and fan capacity. Fuel-rich
operating conditions in the lower furnace region-associated with low NOX
burners can increase the slagging tendency of the coal. The generally longer
flames of low NOX burners will tend to increase furnace exit and superheat/
reheat tube temperatures. Some low NO burners operate with a higher pressure
A
drop or may require slightly higher excess air levels at full load to ensure
good carbon burnout, thus increasing fan requirements.
2-31
-------
Another consideration in retrofitting low NO burners is modifying the
windbox. Modifications may include the addition of dampers and baffles for
better .control of combustion air flow to burner rows and combustion air
distribution to burners within a row. Also, the windbox must be large enough
to accommodate the low NOX burners. If the existing windbox requires
significant modifications to structural components, major re-piping, and/or
windbox replacement, retrofitting low NO burners may not be feasible.
Overfire Air (OFA)--
Overflre air has been installed on approximately 30,000 MW of new
tangentially-fired (T-fired) boiler capacity built since the 1971 NSPS. In
addition an estimated 38,000 MW of pre-NSPS T-fired boilers are believed
technically suitable for OFA retrofit; these units account for an estimated 15
percent of nation-wide pre-NSPS NO emissions (16). OFA can also be used with
wall-fired boilers, but is less effective at controlling NOX emissions than
low-NO burners. As a result, use of OFA is primarily applied to T-fired
boilers.
With retrofit applications, physical obstructions outside of the boiler
may restrict extension of the windbox and make installations of ductwork needed
to supply air to the OFA ports difficult. There must also be adequate distance
between the top burner row, the OFA ports, and the furnace exit for good carbon
burnout and maximum NO reduction performance. Extra fan capacity may also be
necessary.
As in the case of low NOX burners, increased furnace exit gas temperatures
and flame impingement are concerns. To avoid these problems, the distribution
of combustion air between the OFA ports and the burner registers must be
carefully controlled. The Injection velocity required for good OFA mixing with
the combustion gases and the additional pressure drop associated with the OFA
ductwork may require installation of a separate-OFA fan.
2-32
-------
Performance
In new boilers, low NO burners can reduce uncontrolled NO emissions by
" «
40 to 60 percent. Performance in retrofit applications will depend on boiler-
specific considerations, but is generally expected to be somewhat poorer than
in a new boiler.
Overfire air is generally capable of a 15-30 percent NO reduction, with
higher reductions possible in a few cases. OFA can achieve the 1971 NSPS NO
emission limit for wall- and tangentially-fired utility boilers and 1978 NSPS
for NO emission limit for tangentially-fired units.
Cost
Low NO burners and overfire air are relatively inexpensive control
"
methods compared to post-combustion N0x control techniques. NO reduction
performance is the major uncertainty for a retrofit application.
The following table presents the IAPCS estimated capital costs, annualized
costs and cost per ton of pollutant removed for OFA and LNB. The lower capital
and annualized costs are for a large unit with a high capacity factor and the
higher capital and annualized costs are for a small unit with a low capacity
factor. Figures 2.4-1 through 2.4.3 show these costs as a function of boiler
size for overfire air and low NO burners.
Basis for Range Range
Overfire Air -
Capital Cost ($/kW) 1.3 to 6.2
Annualized Cost (mills/kWh) 0.1 to 0.5
Cost Per ton of NO Removed ($/ton) 49.7 to 1,621.2
A
Low NO Burners -
Capital Cost ($/kW) 5.5 to 25.5
Annualized Cost (mills/kWh) 0.2 to 2.1
Cost Per ton of N0y Removed ($/ton) 79.9 to 2,388.8
^
Appendix B contains tables for estimating the cost of combustion controls as a
function of boiler size, capacity factor and NO reduction.
2-33
-------
oc
in
a.
OT
Legend
• Overfire Air
+ Lo* Nox Burners
Copocity Foctor = 50%
100
300
MW
700
1000
1300
Figure 2.4-1. NOx Combustion Controls - Capital Cost,
Current 1988 Dollars
1.2
1.1 -
1 -
0.9 -
0.8 -
0.7 -
0.6 -
0.5 -
0.4 -
0.3
0.2 H
0.1
Legend
• Overfire Air
•*• Low Nox Burners
Capacity Factor = 50%
100
300
500
700
1000
1300
MW
Figure 2.4-2. NOx Combustion Controls - Levellzed
Annual Cost, Current 1988 Dollars
2-34
-------
If
fc
Overfire Air - 10% NOx Reduction .
Overfire Air - 30% NOx Reduction
Low NOx Burners - 40% NOx Reduction
Low NOx Burners - 55% NOx Reduction
Copocity Factor- = 50%
TOO
300
500
700
1000
1300
MW
Figure 2.4-3. NOx Combustion Controls - Cost Per Ton
of NOx Removed, Current 1988 Dollars
2-35
-------
2.5 THROWAWAY WET FGD
Throwaway wet FGD technologies are the most widely used method of post-
combustion SO. control in the U. S. This section discusses the two major
processes used in the U. S.: lime/limestone scrubbing and dual alkali The
use of additives to enhance lime/limestone performance is also discussed.
Additional processes are in use 1n Japan and Europe and have been
pilot-scale tested in the U. S. These processes Include Chiyoda Thorough-
bred 121 and Dowa from Japan and Saarberg-Holter from West Germany. For
information on these processes refer to Reference 18.
Description
Lime/Limestone Scrubbing--
Lime/limestone scrubbing is the most common post-combustion S0? control
technique currently applied to utility boilers. A process flow diagram is
shown in Figure 2.5-1. There are four fundamental steps to the process:
preparation of the Hme or limestone slurry, contacting the flue gas and
slurry, reaction of the lime or limestone with S0_, and removal of solid waste.
Reagent preparation involves crushing limestone or slaking lime to produce
a slurry of fine hydrated lime or limestone particles. Although many existing
systems use lime, most future systems are expected to use limestone due to its
lower cost. Flyash from alkaline western coals has also been used in
conjunction with lime and limestone. The slurry is fed to the scrubbing loop
which consists of an absorber (usually a spray tower) and reaction tank; the
scrubber system for most utility plants consists of multiple absorber/reaction
tank modules. SO. 1s removed from the flue gas (which has already been cleaned
of flyash by an ESP or fabric filter) In the absorber while the reaction tank
provides time for limestone dissolution and reaction with the dissolved S0?.
The product of this reaction is a mixture of crystalline calcium sulfate and
calcium sulfite. Small droplets of slurry which become entrained in the flue
gas are captured by a mist eliminator. Reacted solids are removed and
dewatered with the clear liquor being returned to the process. The solid waste
is either ponded or landfilled.
2-36
-------
Stack Gas
Flue Gas
Organic
Acid Addition
(Optional)
Reagent
Preparation
Lime or
Limestone
Absorber
Slurry
1
Reaction
Tank
Slurry
Clear Liquor
Solids
Preparation
T
Waste
Disposal
Figure 2.5-1. Lime / limestone FGD system flow diagram.
2-37
-------
During the initial applications of lime/limestone FGO systems a variety of
chemistry problems were encountered, often with the result being rock-like
gypsum (calcium sulfate) scale formation on equipment. These problems have
been overcome through experience and better understanding of the complex
chemical interactions. Today, most lime/limestone FGD problems involve
adjusting the chemistry to optimize raw material utilization and SO* removal.
However, materials selection continues to be a problem. To reduce impacts on
boiler operation, scrubber systems are generally designed with an extra
absorber module which can be operated when any of the other modules are down
for maintenance or repair.
One technique for Improving the performance of limestone FGD systems is
the use of organic acid additives, such as adipic and other dibasic acids.
This improves SO. removal and enhances other aspects of the process, but does
not greatly increase operating costs. As a result of EPA Air and Energy
Engineering Research Laboratory (EPA-AEERL) sponsorship of development efforts,
this technique has been successfully used in retrofit applications at several
utility Installations. Use of organic acids allows existing FGD systems to
improve system reliability and SO- removal without affecting other aspects of
the process or requiring significant additional S0» control equipment.
Dual Alkali Scrubbing--
Dual alkali scrubbing was developed to avoid the problems of erosion,
scaling, and solids deposition found with lime/limestone FGD. As the name
implies, two alkalis are used: a clear solution (generally sodium sulfite) in
the absorber followed by addition of lime in the reaction tank to regenerate
the spent solution for recycle to. the absorber. The resulting calcium
sulfite/sulfate sludge is then dewatered and landfilled. Makeup sodium
(typically soda ash) Is added to the regenerated solution to replace residual
sodium lost in the filter cake.
Several variations of the dual alkali process have been investigated
(e.g., using ammonia and potassium as the first-stage alkali). However, the
use of sodium is the most advanced and is the only process in commercial use in
the U. S. A simplified diagram of the sodium/lime process is shown in
Figure 2.5-2.
2-38
-------
Lime silo
Flue gas
by-pass
Scrubbed gas to
upper stack breechlngs
Clean gas to lower
stack breechlngs
CA(OH)2
Flue gas
from
electrostatic
preclpltator
Soda Ash
CaO
Water
Figure 2.5-2. Simplified flow diagram for a dual akali FGD system
Source: Reference. 1.
Solid to on-slte
landfill
-------
Sodium-based, dual alkali systems are more suitable for high-sulfur than
for low-sulfur (less than 1.5 percent) coal applications. With low-
sulfur coals, a greater percentage of the sulfHe is oxidized to sulfate. This
lowers the available alkalinity in the scrubber and, thus, lowers S02 removal
capacity.
Additive-Enhanced FGD--
Several additives have been added to lime/limestone FGD systems to improve
SO. removal, sorbent utilization, and/or system operation. These include:
adiplc acid, mixtures of dibasic organic acids (DBA), magnesium, and
thiosulfate. The decision to use a specific additive depends upon system
chemistry and performance. In general, except for the additive feed system,
there are no major hardware differences compared to a conventional lime/
limestone system.
Organic acids such as adipic or DBA are currently the most widely used
additives. These acids buffer the gas/liquid contacting interface at a pH of 3
to 5, thus improving the driving force for S02 removal. .For example, addition
of only 2000 ppm of adipic acid can more than double the normal dissolved
alkalinity of the absorbing limestone slurry, and thus can compensate for
inadequate gas/liquid contact in the absorber. Organic acids buffering also
allows operation at lower pH than limestone-only systems, without loss of SO.
removal capability. This results in higher limestone utilization which can, in
some cases, reduce mist eliminator scaling problems.
Magnesium-promoted FGD involves addition of magnesium-containing (i.e.
dolomitic) lime or magnesium sulfate (MgSO.) to either a lime or limestone
system. Magnesium salts are more soluble than the corresponding calcium salts,
providing higher liquid-phase alkalinity and increased absorption of SO..
Sodium thiosulfate (Na.S.O.) addition has been used to reduce scaling in
both lime and limestone systems and to improve sludge dewatering. Sodium
thiosulfate is also a very active inhibitor of sulflte oxidation; as little as
100 ppm S203 can virtually halt oxidation of calcium sulfite to calcium
sulfate. However, consumption of this additive can be high due to degradation
2-40
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of thiosulfate to other species, such as trithionate, which do not inhibit
oxidation. This approach was commerciany demonstrated in 1985 at Seminole
Electric Cooperative's plant near Palatka, Florida.
Applicability
Lime, limestone, sodium carbonate, dual alkali, and additive-enhanced FGD
are commercially available and have been widely applied to utility boilers.
The total capacity of operating and contracted throwaway FGD systems as of
December 1985 exceeds 68,000 MW as shown in Table 2.5-1. There is currently an
estimated 4,000 MW of scrubber capacity using additives (19). From a process
viewpoint, lime/limestone FGD is applicable to all boilers. Dual alkali
systems are best suited to boilers burning high-sulfur and high-chlorine coals.
Depending upon the arrangement of the boiler and associated equipment, a
scrubber retrofit may be simple or difficult. Difficult retrofits often
require either long runs or complex arrangements of ductwork, which can be very
expensive and time consuming to install. In addition to the absorber and
reaction tank in which SO* removal actually occurs, additional space is also
required for sorbent receiving and storage, slurry preparation, dewatering
of reacted sorbent, and solid waste disposal. Replacement or relining of the
stack may also be required to eliminate corrosion problems caused by the wet
flue gas. Modification of the boiler may also be required to balance and
control pressure changes caused by retrofitting a scrubber and additional
fans. The high capital cost of the equipment combined with retrofit
difficulties makes it very expensive to apply lime/limestone FGD on plants
operating at low capacity factors or with little remaining plant life.
In the U. S., additive enhancement is currently used solely as a retrofit
method for upgrading the performance of existing FGD systems. The minimal
changes in system design required for addition of additives make retrofit
relatively easy. Future systems may be designed with additives to reduce
capital and operating costs and to boost FGD removal efficencies well above
90%.
2-41
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TABLE 2.5-1. TOTAL OPERATING AND PLANNED THROWAWAY FGD CAPACITY
BY U. S. ELECTRIC UTILITIES (as of December 1985)
Units
Capacity Factor
Limestone
In operation
Under construction
Contracts awarded
Lime
60
9
_8
77
26,008
17,113
1.505
36,977
In operation
Under construction
Contracts awarded
Sodium Carbonate
In operation
Under construction
Contracts awarded
Dual Alkali
In operation
Under construction
Contracts awarded
Total
6
1
_2
9
150
17,113
0
2.036
19,149
1,505
550
1.100
3,155
1,963
265
0
2.228
68,072
Source: Reference 19.
2-42
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Performance
Performance of FGD systems is measured with respect to both SO- removal
efficiency and system availability. In general, both lime/limestone and
dual alkali are capable of up to 95 percent SCL removal with careful design and
operation. Removals of 90 percent are more common. As discussed earlier,
additives such as adipic acid frequently double the dissolved alkalinity of
limestone slurries. This improvement in alkalinity can reduce SO- emissions
from existing scrubbers by a factor of two or more while at the same time
increasing unit availability.
In general, lime/limestone scrubbers tend to be more chemically complex
in plants using high-sulfur coals versus low-sulfur coals, and thus availabi-
lity is lower. Addition of spare absorber modules can improve overall system
availability by providing backup units and allowing more time for maintenance
and repair of each module. However, new scrubber design and the use of
additives reduce the need for module sparing in the future.
Cost
The range of estimated costs using the IAPCS cost model for both new and
retrofit lime/limestone FGD systems applied to coal-fired utility boilers are
presented below. Depending on local sorbent availability, waste disposal
options, and other site-specific factors, these costs may vary somewhat, but
will generally be within the range shown. The lower capital and annualized
costs are for large units with high capacity factors, low retrofit factors and
low sulfur coal content, while the higher capital and annualized costs are for
small units with low capacity factors, high retrofit factors and high sulfur
coal content. Appendix C contains tables of costs as a function of boiler
size, capacity factor, coal sulfur and retrofit difficulty.
Basis for Range Range
Capital Cost ($/kW) 125 to 628.5
Annualized Cost (mills/kWh) 11 to 88.6
Cost Per ton of S02 Removed ($/ton) 508.3 to 10,618.6
These costs represent conventional design practices for new installations
relative to the size of absorbers and installation of backup systems to assure
2-43
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compliance with New Source Performance Standard (NSPS) reliability
requirements. If these design practices are altered to comply with less
stringent reliability requirements, larger—but fewer—absorbers might be
installed. Analysis of cost savings at several plants using EPA's IAPCS
computerized FGD costing model found this altered design philosophy could
reduce capital cost by 16-32 percent and total annual costs by 9-12 percent
relative to conventional designs. Under the Department of Energy Clean Coal
program, the wet limestone Mitsubishi FGD process was selected for commerical
demonstration. The program will demonstrate the technical and economic
feasibility of an advanced, single, 600 MW absorber system to obtain 90-95% SOp
removal on high sulfur (>3 % S) Indiana/Illinois coals.
The costs and benefits from additive enhancement depends in large part on
the performance of the existing FGD system. For a limestone FGD unit operating
at 90 percent SCL removal, but at high liquid-to-gas absorber ratios, use of
organic acids can significantly reduce operating (i.e., slurry pumping) costs
while maintaining the same SO. removal; EPA's IAPCS model estimates a total
cost savings of roughly six percent at these conditions. For a unit needing to
increase its SO- removal efficiency, the economic advantages of adding organic
acids may be significantly higher. In general, the capital costs of converting
an existing limestone FGD system to use these additives are minimal and quickly
pay for themselves if improved system reliability is achieved.
Engineering estimates of capital and fixed O&M costs for dual alkali
systems are generally lower than for lime/limestone systems (1). However,
total costs (including variable O&M) are similar due to the high cost of lime.
Enhancement of the dual alkali process to allow use of limestone rather than
lime for regenerating the scrubber solution could significantly reduce the cost
of the dual alkali process. Detailed calculations of dual alkali costs are not
included in this report.
The costs for a specific FGD system are a function of many variables
including unit size, coal sulfur level, percent SO. removal, and load factor.
As previously mentioned, the remaining llfespan of a plant is also important.
An operator does not want to make large capital Investments in a plant that has
a short remaining lifetime. Also, smaller plants (less than 100 MW) are much
more expensive per kW to retrofit with scrubbers than larger plants.
Figure 2.5-3 through 2.5-5 present costs as a function of plant size and coal
sulfur content.
2-44
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O
O.
to
C
UJ
QL
(A
Legend
• 1% s
+ 2% S
o 3% S
& 4% S
Capacity Factor = 50%
Limestone Cost = $l7.3/ton
Retrofit Factor = 1.0
100
300
500
MW
700
900
1,100
1,300
Figure 2.5-3. L/LS FGD - Capital Cost, Current 1988 Dollars
Legend
+
o
1% s
2% S
3% S
4% S
Copacity Factor = 50%
Limestone Cost = $17.3/ton
Retrofit Factor = 1.0
12
100
300
500
700
900
1,100.
1.300
MW
Figure 2.5-4. L/LS FGD - Levelized Annual Cost,
Current 1988 Dollars
2-45
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0.5
100
300
Legend
• 1% s
+ • 2% S
O 3% S
a 4% 5
Copacity Factor = 50%
Limestone Cost = $17.3/ton
Retrofit Factor = 1.0
1.100
1,300
Figure 2.5-5. L/LS FGD - Cost Per Ton of S02 Removed,
Current 1988 Dollars
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2.6 BY-PRODUCT RECOVERY FGD TECHNOLOGIES
Description
Recovery of sulfur from SO- as a useful product rather than discarding it
is beneficial with respect to resource conservation, and also because of the
potential improvement in process economics. Numerous processes have been
proposed in an effort to achieve these goals. These processes vary widely in
sorbent used, product made, and stage of development. Currently, three types
of processes have attained commercial status.
t Gypsum recovery - Conventional FGD system with forced oxidation
and additional process features to produce salable gypsum.
t Wellman-Lord - Reaction of sodium sulflte with S02 to form sodium
bisulfite followed by steam stripping to recover sulfuric acid or
elemental sulfur.
t Magnesia slurry - Reaction of MgO/Mg(OH)- with S02 to form magnesia
sulfite which is then separated, dewatered, dried, and calcined to
recover sulfuric acid.
Gypsum Recovery--
In general, wet lime/limestone FGD requires several process enhancements
to supply consistently pure gypsum. These enhancements include forced
oxidation to produce the required gypsum content in the slurry solids, particle
agglomeration, and washing of gypsum filter cake. The gypsum slurry is then
dewatered using conventional equipment, resulting in a cake that is greater
than 80 weight percent solids (20, 21). Approximately 2.7 tons of gypsum is
produced per ton of S02 removed from the flue gas (22). Washing of product
gypsum, especially with high-chloride coals, generates a liquid waste stream.
Wellman-Lord--
The main features of the Wellman-Lord process (Figure 2.6-1) are the use
of a sodium-based sorbent, regeneration of the sorbent solution, and conversion
of S02 to sulfuric acid or elemental sulfur. The scrubbing portion
2-47
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Doubte-Loop
Absorber
oo
S(X Rich Stream
lo Elementate or
SuHuric Acid
Low
• Pressure
Steam
Sodium SulfalB
Solids Purge
Figure 2.6-1. Wellman-Lord Process Schematic
-------
of the process involves circulating sodium sulfite solution, typically in a
plate-type absorber, to absorb SO.. Sodium bisulfite is the primary reaction
product, though some oxidation to sodium sulfate will occur depending on the
flue gas oxygen content. This sulfate is purged as a solid, using a
crystallizer followed by drying. The sodium bisulfite solution from the
absorber is thermally decomposed in a steam-heated evaporator to regenerate
sodium sulfite. Solids formed due to evaporation are redissolved in a separate
vessel; at this point make-up soda ash (Na^CO,) is added as necessary.
Overhead vapor from the evaporator is an SOg-rich stream, suitable for
conversion to elemental sulfur or sulfuric acid using commercial technology. A
purge stream requiring treatment and disposal is produced in order to keep
thiosulfate levels (i.e., dissolved solids) below a threshold.
Magnesia Slurry--
In the magnesia slurry process (Figure 2.6-2), an aqueous slurry of
regenerable magnesium hydroxide and magnesium sulfite is used to absorb SO.
from flue gas. Reaction products formed in the absorber include several
magnesium sulfite hydrates and magnesium bisulfite. A portion of the reacted
slurry from the main scrubbing loop is sent to a thickener for partial
dewatering. A centrifuge is then used to recover a wet cake of magnesium
sulflte/sulfate crystals. The crystals are sent to a rotary or fluidized bed
dryer while the liquor is returned to the circulation loop. Following the
solids drying step, decomposition of the magnesium sulfite/sulfate compounds is
carried out in a high-temperature calciner. This liberates a concentrated SO.
gas stream that can be routed to the sulfur (or sulfuric acid) production unit.
The calcined solids are regenerated as MgO for recycle to the absorber loop.
Applicability
Gypsum Recovery--
Gypsum recovery technology can be used almost anywhere ordinary lime/
limestone scrubbing is appropriate. It is only necessary to add forced
oxidation, product solids washing (to assure adequate gypsum purity), and
2-49
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Rehealar
SO, Absorber
FhjeGas—»
To Slack
Coke
Catelner
Figure 2. 6-2. Magnesia Slurry Process Schematic (23)
-------
shipping facilities. The primary commercial use of this technology is to
produce gypsum for wall board or as a cement retarder. Approximately 20 forced
oxidation units are operating in the U. S. However, only two plants have plans
to recover gypsum for sale. In Japan, FGD gypsum is routinely substituted for
natural gypsum, and practically all the lime/limestone FGD installations in
Europe successfully market recovered gypsum. The main motivation for
by-product utilization overseas is lack of space and high cost for waste
disposal.
Annual consumption of raw gypsum in the U. S. was approximately 20 million
tons as of 1984. The wallboard industry consumes about 70 percent of this
amount, with the balance used in the manufacture of cement and plaster, and in
agriculture (22). Wallboard manufacturers historically considered only purity
and grain-size; however, for recovered FGD gypsum other factors must be
considered. For example, as soluble salt concentration increases the gypsum
calcination temperature decreases, which may result in weakening of the bond
between the paper coating and gypsum core.
Wellman-Lord--
The Wellman-Lord process has been retrofit on coal-fired utility boilers
in both the United States and Japan. Important features of the technology
include SO- removal levels in excess of 90 percent over a wide range of inlet
SO2 concentrations (24) and minimal scaling and plugging. However, as with
other regenerate FGD processes, the Wellman-Lord system is mechanically
and chemically complex. Also, handling the concentrated SO^ product stream as
elemental sulfur or sulfuric acid entails additional complex processing, and
advantageous economics require an available market for the end product.
Another limiting factor for utility applications may be the disposal of the
sodium sulfate solids, although the waste volumes are considerably less than
for throwaway FGD systems.
The Wellman-Lord system is commercially available and as of 1985 has been
applied to the four U. S. utility plants listed in Table 2.6-1 (25).
2-51
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TABLE 2.6-1. WELLMAN-LORD UTILITY INSTALLATIONS IN THE U.S. (24)
Utility Plant Name (MW) %S Fuel SC>2 End Product Start-up
N. Indiana
Public Ser.
Public Ser.
Co. of N. M.
Oelmarva
Power & Lt.
Public Ser.
Co. of N. M.
Mitchell (115) 3.0
San Juan (700) 0.8
Delaware City 7.0
(180) (coke)
San Juan (1100) 0.8
Elemental S
Elemental S
Sulfuric acid
Sulfuric acid
1976
1978
1979
1982
2-52
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These applications represent both low sulfur (0.8 percent) and high sulfur (7
percent) coals. In early 1986, It was reported that one German and one
Italian utility had ordered construction of Wellman-Lord units; the latter
being a 10 MW demonstration plant (26). In Japan there are two utility
Installations, one at Chubu Electric's Nishinagoya station (620 MW) and the
other at the Niigata station (380 MW) of Tohoku Electric (1).
Retrofits are possible, but there must be adequate land available near the
boiler and It must be possible to modify the existing ductwork to route the
flue gas to and from the FGD system. Many of the Industrial sites where the
Wellman-Lord system is employed are existing sulfurlc add or sulfur plants.
Magnesia Slurry--
The magnesia slurry process has been demonstrated on both oil- and
coal-fired utility systems. Retrofit is possible and the process has good
turn-down capability. The technology is well developed; in the mid-1970's two
demonstration units having capacities of 90 and 150 MW operated for
2 years (27). Three commercial units totaling 724 MW were brought into service
by Philadelphia Electric Company in 1983 (1). MgSO, is shipped off-site to
existing sulfuric acid plants and the regenerated MgO is returned to the power
plant site.
Performance
S02 removal performance of gypsum by-product recovery plants is about the
same as conventional lime/limestone FGD systems. However, system operation
must be more carefully controlled in order to avoid excessive unreacted reagent
in the product solids. The 2250 MW Martin Lake plant has operated since late
1984 at better than 90 percent SO- removal, with greater than 95 percent
oxidation and 90 percent utilization of limestone reagent. The plant burns a
lignite with a sulfur content of 1-4 percent. The highly oxidized sludge
clarifies quickly, but careful solids handling is needed to avoid line
pluggage. Also, close pH control is necessary to maintain a sufficiently high
oxidation level.
2-53
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The 160 MW Muscatlne, Iowa plant has successfully operated a forced
oxidation system at 94 percent sulfur removal with 3.5 percent sulfur coal
(20). After initial start-up in early 1984, the operability of the system has
averaged better than 95 percent, and reliability above 99 percent. There have
been some problems with scaling of gypsum on the scrubber internal components
due to the high oxidation level. The largest problem to date with the produced
gypsum has been excess unreacted limestone remaining in the solids.
Commercial Wellman-Lord systems have demonstrated removal of SO- in excess
of 90 percent. High SCL removal is achievable due to the high alkalinity of
the sodium-based sorbent solution. In addition, the process can handle widely
varying inlet SCL concentrations, and scaling or solids-plugging do not present
a problem. However, process complexity has resulted in greater maintenance
requirements than for conventional lime/limestone FGD operation.
Cost
The costs for producing salable gypsum include many components, but are
partially offset by disposal cost savings. In general, the additional
operating costs for gypsum production roughly balance those for disposal The
key. variable then becomes whether the market: and price for the product can
justify the additional capital costs of increased washing capacity, forced
oxidation, or other processing. Another factor affecting this option in the
U.S., was a tax depletion allowance of 14 percent (as of 1984) granted to
gypsum mining companies (22). This tax credit reduces the actual cost of mined
gypsum, making it more difficult for FGO gypsum to be competitive.
Table 2.6-2 presents estimated capital costs, operating costs, and cost
effectiveness for a Wellman-Lord FGD system as applied to a coal-fired utility
boiler using EPRI guidelines (28). The Wellman-Lord process is more costly
than conventional lime/ limestone FGO systems. However, the capital cost
differential of the two systems decreases with lower sulfur coals. Operating
costs are considerably higher for the Wellman-Lord process, as might be
2-54
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TABLE 2.6-2. COST ESTIMATES FOR WELLMAN-LORD FGD
(Current 1988 $)
Technology
New 500-MU Basel oad Power
Capital Cost, S/kW
M1lls/kWh, Constant $
Current $
$/Ton, Constant $
Current $
% SO-
Removed
Plant:
90
90
90
90
90
3.5%
259 -
11.7 -
20.0 -
443 -
763 -
S
290
12.1
21.0
461
795
2%
204 -
9.4 -
16.1 -
624 -
1,070 -
S
229
9.7
16.7
647
1,21
Retrofit 500-MW Baseload Power Plant1:
Capital Cost, $/kW 90 335 - 375 265 - 296
Mills/kWh, Constant $
Current $
$/Ton, Constant $
Current $
Retrofit 250-MW Intermediate
Capital Cost, $/kW
Mills/kWh, Constant $
Current $
$/Ton, Constant $
Current $
90
90
90
90
Load Power
90
90
90
90
90
12.8
22.1
484
832
Plant2:
538
31.5
44.1
1,197
1,679
- 13.3
- 23.2
- 506
- 870
- 602
- 33.8
- 47.5
- 1,289
- 1,811
10.2 -
17.7 -
680 -
1,169 -
424 -
25.7 -
35.9 -
1,712 -
2,398 -
10.6
18.6
709
1,270
475
27.6
38.7
1,839
2,579
Retrofit difficulty of 1.3 and capacity factor of 65.
Retrofit difficulty of 1.6 and capacity factor of 35.
2-55
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expected in view of the greater mechanical and chemical complexity. Retrofit
costs are also substantially higher relative to new Installation of equivalent
or greater capacity lime/limestone systems. In these cost comparisons, the
market potential for recovered by-product at a particular-location is not
expllclty included.
Extremely limited experience, to date, with the magnesia slurry process on
a commercial scale would make specific cost comparisons unreliable. However,
it is generally recognized that capital and O&M costs of the magnesia process
are higher than conventional lime/limestone FGD.
2-56
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2.7 SPRAY DRYING
Description
Spray drying is an established technology for desulfurization of flue gas
from high- and low-sulfur coals. The first installation of utility spray dryer
systems occurred in the late 1970's. Thirteen systems, representing nearly
4,300 MW of generating capacity, are currently in service, and an additional
five systems representing over 2,800 MW of generating capacity are in design or
construction phases (29). All of these full-scale systems are based on coals
containing less than two percent sulfur. Evaluation of spray dryer systems for
utility boilers burning greater than three percent sulfur coal was initiated in
1985 using two 1 MW, pilot-scale spray dryers at TVA's Shawnee Test Facility.
A typical spray dryer system is shown in Figure 2.7-1. Lime is the
reagent in 17 of the utility systems mentioned above, while one (the first
spray dryer system purchased) uses a sodium-based reagent. Only lime-
based systems are described here.
In a lime-based system, quicklime is slaked to form calcium hydroxide, and
the resulting slurry is combined with additional process makeup water and (in
most systems) recycled solids. This combined slurry is then atomized and mixed
with the flue gas at air heater exit conditions in the spray dryer vessel.
Slurry atomization is generally accomplished with a rotary atomizer. Two-fluid
air atomizing nozzles are used in some systems. Several simultaneous reactions
occur immediately following atomization:
• Water in the slurry droplet evaporates;
• The evaporating water cools and humidifies the flue gas; and
• Flue gas SO- reacts with the lime, producing calcium sulfite
and calcium sulfate solids.
The amount of water fed to the spray dryer is carefully controlled to avoid
complete saturation of the flue gas. However, a close approach to
2-57
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Head Tank
Lime
Feed
Bin
Flue Gas from Unit 7
Dilution
[ritlon Water
ker
in
CD
Slaking
Water
Recycle
Storage
Silo
Fabric Fitter
lime Slurry
Storage Trough
Mix Tank
Atomizer
Feed Tank
I.D. Fan
\ /
c
V.
1
Recycle Solids
Y
Solids to
Disposal
Stack
Figure 2.7-1. Lime Spray Drying Process Flow Diagram (30).
-------
saturation temperature (20°-35°F) is needed to achieve high SO- removal and
lime utilization. Sorbent feed rate, degree of slurry atomization (i.e.,
droplet size), flue gas residence time in the spray dryer vessel, and the
approach to flue gas saturation temperature are carefully controlled to achieve
acceptable S02 removal performance while avoiding problems with drop out or
caking of incompletely dried solids in the spray dryer vessel. The calcium
content in the slurry is limited to a maximum of about 35 percent solids.
Typical flue gas residence time in the spray dryer vessel is 8-12 seconds.
The nearly-saturated flue gas leaves the spray dryer vessel and enters the
particulate collection device which is either a fabric filter or an
electrostatic precipitator (ESP). Most utility experience to date has been
with fabric filtration. In the particulate collection device, the original
coal flyash and the dried sorbent are removed from the flue gas. Additional
SO- removal occurs across the particulate control device. Fabric filters
contribute as much as 20-30 percent to overall system S02 removal (30). Less
information is available on SOL removal across an ESP, but it may be as high as
25 percent at low approach-to-saturation temperatures (20°F) (31).
In most systems, a portion of the collected solids are recycled to the
slurry mix tank. Solids recycle offers several advantages over once-through
lime operation, including improved lime utilization, improved rotary atomizer
operation, faster droplet drying, and reduced solid waste (32).
Lime spray drying has several perceived advantages over wet lime/limestone
scrubbing. First, the need to recirculate large quantities of slurry are
avoided. L1quid-to-gas volume ratios for spray dryer systems tend to be 10 to
50 times smaller than what is typical for lime/limestone scrubbing. Second,
wastes from the spray drying process are produced as dry solids, rather than as
a wet sludge. However, moving and storing relatively large quantities of dry
solids can also be a troublesome operation. Third, fewer corrosion problems
occur with Hme spray drying since the system is dry rather than wet.
2-59
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Applicability
Although spray drying 1s generally considered a commercially established
technology for low-sulfur coal utility applications, data with high-sulfur
coals are limited. In high-sulfur coal applications, heat balance
considerations will limit SO- removal by limiting the amount of lime and water
that can be added as well as the amount of sorbent recycle. Specifically, the
differential between the spray dryer inlet temperature and the approach to flue
gas saturation temperature at the spray dryer outlet will determine the amount
of water that can be evaporated by the flue gas. If too much water is used,
the flue gas will cool to below saturation temperature and result in scaling
and corrosion problems. If too little water is used, poor slurry atomization
and low lime utilization will occur. Additives (specifically chlorides) have
been successfully used in short-term tests to Increase sorbent utilization at
high SO- removal rates.
Spray dryers can be used in retrofit applications 1f there 1s sufficient
land near the boiler and If ductwork can be cost-effectively run from the air
heater outlet to the spray dryer inlet and from the spray dryer outlet to the
particulate control device.
The type of existing particulate control device may influence the
applicability of spray drying as a retrofit FGO technique, particularly for
high-sulfur coal applications. Fabric filters are well suited for spray dryer
retrofits, but few existing utility boilers are so equipped (particularly in
high-sulfur coal service). Existing ESPs have several limitations, including:
• Less contribution to system SCL removal than a fabric filter;
• Greater sensitivity to increased particulate loading caused by the
sorbent (addition of a spray dryer will require greater particulate
removal efficiencies to maintain emission rates at pre-spray drying
levels); and
• Potential corrosion problems.
However, these Impacts are poorly quantified and more testing is needed.
2-60
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For low-sulfur coal applications where high recycle rates are achievable
(30), adequate S(L removal performance may result even with relatively low SO-
removal in the particulate control device. Also, most existing ESPs designed
for low-sulfur coal have high specific collection area (SCA) values of 400
2
ft /thousand acfm and greater. With high initial SCA values, the beneficial
effects of gas volume shrinkage and gas conditioning which results from cooling
and humidification in the spray dryer may more than offset the effects of
increased mass loadings.
Many existing ESPs in high-sulfur coal applications have SCA values of 200
or less. At these SCA values, it is unlikely that existing emission rates will
be maintained after a spray dryer is installed upstream. Pilot studies
conducted by TVA (31) indicate that an ESP with an SCA of 300 which currently
meets existing particulate emission regulations by a comfortable margin should
be able to maintain compliance after a spray dryer is added. If the ESP cannot
be adequately upgraded, installation of a replacement baghouse would be
required.
Other issues for spray dryer retrofits include increased corrosion of
particulate control devices due to operation at lower gas temperatures and
higher moisture content, and the need to upgrade existing particulate collector
solids handling equipment to deal with increased loadings and alkalinity caused
by the spray dryer. Replacement of wet sluice ash handling systems with a dry
system would also be required to preheat plugging of the sluice lines.
Performance
All existing and currently planned lime spray dryer FGD systems are
designed for coals with less than two percent sulfur; design S02 removals are
61 to 90 percent (29). Most of the 13 operating units readily meet their
design S02 removal levels. Less experience is available for high-sulfur coal
applications. Long- term utility test results are not yet available.
Short-term spray dryer tests combined with fabric filtration have been
conducted on high-sulfur coals at Argonne National Laboratory, General Motors'
Buick Steam Plant, and Northern States Power Company's Riverside Station.
These tests demonstrated over 90 percent SO, removal on coals up to 4.2 and 3.2
-» . £ '
percent sulfur with and without the use of additives, respectively (33). In
2-61
-------
pilot-scale testing using a lime spray dryer/baghouse in high-sulfur coal
applications, TVA demonstrated that greater than 80 percent removal is
feasible. Lime utilization at these SO- removal rates was 60-65 percent.
Pilot-scale tests for a high-sulfur spray dryer/ESP system conducted by
Wheelabrator A1r Pollution Control and the U.S. Department of Energy achieved
SOp removal levels as high as 90 percent with a 2.7 percent sulfur coal (34).
Lime utilization was only about 30 percent; however, these tests did not employ
recycle and the approach to flue gas saturation was not as close as is possible
to maximize lime utilization. The spray dryer inlet temperature for these
tests was rather high at 400°F.
Tests conducted by TVA on a 10 MW spray dryer/ESP pilot unit at Shawnee
with a 3.5 to 4 percent sulfur coal and a 300°F inlet flue gas temperature
achieved overall SO- removal levels of 90 percent (35). Most of these tests
were conducted at a 18°F approach to adiabatic saturation at the spray dryer
outlet.
The data discussed above indicate that high levels of S02 removal
(90 percent) are possible for new units that are designed to operate at these
levels. However, for retrofit applications where the existing ESP will be
reused, SO. removal may be limited by ESP performance. EPRI also is sponsoring
research on lime spray drying at Its High Sulfur Test Center. Pilot-scale
tests on a 4 MW' spray dryer with fabric filter have achieved 90% sulfur dioxide
removal on coals with 3.5-4% sulfur and a calclum-to-sulfur ratio of 1.3 to
1.6.
The potential for combined SO./NO reduction also exists by addition of
roughly 10 weight percent sodium hydroxide to the lime. The NaOH both reacts
with SCL and extends particle reactivity as well as serves as a catalyst to
reduce NO emissions. In pilot-scale tests with a three percent sulfur coal
and a lime-to-sulfur ratio of 1.3, SO- and NO reductions of 95 and 55 percent,
respectively, were achieved. However, the process oxidizes a portion of the NO
to NO-, resulting in increased plume opacity and discoloration (36). Ammonia
injection has been demonstrated to reduce the NO. plume.
2-62
-------
Cost
Due to the sensitivity of spray dryer performance to plant design and
operating variables, capital and operating costs for retrofit spray dryer
applications are expected to be more site specific than wet lime/limestone FGD.
Key factors in evaluating spray dryer costs at a given site include air heater
outlet temperature, length of ducting between the air heater outlet and spray
dryer inlet, percent sulfur in the coal, and whether the existing particulate
control device (typically an ESP) can be upgraded to handle the increased
particulate loading resulting from spray dryer operation. The expected SO-
reduction across a baghouse versus an ESP is also important..
A critical consideration in spray dryer economics will be whether to use
the existing ESP or install a new baghouse. Installation of a new baghouse can
increase retrofit capital costs by 30-50 percent and annualized costs by 25-40
percent above the cost of a spray dryer system in which the existing ESP can be
used. Because of the uncertainty in spray dryer performance and cost, two cost
cases are examined. One case assumes that the existing ESP can be upgraded and
continues in service. The other assumes that a new baghouse will be installed.
The following table presents the range of values estimated using IAPCS
cost model for capital cost, annualized cost, and cost per ton of pollutant
removed. The lower costs are for a large unit with a low retrofit factor
burning low sulfur coal. The higher costs are for a small unit with a high
retrofit factor burning high sulfur coal. Figures 2.7-2 through 2.7-4 show
these costs as a function of boiler size, coal sulfur content, and particulate
control option.
2-63
-------
o
a.
JC
u
o
a
Lime Cost = $63.3/ton
ESP-SCA = 400 fr/lOOO ACFM
100
100
300
500
700
MW
(a) Existing ESP
900
i. 1 DO
1,300
1% S •
2% S
3% ^
4% S
Capacity Factor = 50/5
Lime Cost = $63.3/lon
ESP-SCA = 400 ft2/1000 ACFM
300
500
MW
700
900
1.100
1.300
(b) New Baghouse
Figure 2.7-2. Lime Spray Drying - Capital Cost,
Current 1988 Dollars
2-64
-------
tc
\u
OL
QC
UJ
0.
Capacity Factor = 50%
Lime Cost = $63.3/ton
ESP-SCA = 400 reViooo ACFM
100
TOO
300
500
MW
700
900
1100
1300
(a) Existing ESP
1% S
2% S
3% S
4% S
Capacity Factor = 50%
Lime Cost = $63.3/ton
ESP-SCA = 400 ft2/!000 ACFM
300
500
MW
700
900
1100
1300
(b) New Baghouse
Figure 2.7-3. Lime Spray Drying - Annualized Cost,
Current 1988 Dollars
2-65
-------
CT3
OC
*•«
l. <0
03
Q.O
-
Capacity Factor = 50%
Lime Cost = $63.3/ton
ESP-SCA = 400 ftVlOOO ACFM
100
100
300
500
MW
•700
900
1100
1300
(a) Existing ESP
1% s
2% S
3% 5
4% 5
Capacity Factor = 50%
Lime Cost = $63.3/ton
ESP-SCA = 400 ftVlOOO ACFM
300
500
700
900
1100
1300
MW
(b) New Baghouse
Figure 2.7-4. Lime Spray Drying - Cost Per Ton of S02 Removed,
Current 1988 Dollars
2-66
-------
Range
Existing ESP -
Capital Cost ($/kW) 64.7 to 535.4
Annualized Cost (mills/kWh) 6.3 to 65.7
Cost Per ton of SC>2 Removed ($/ton) 441.5 to 8,084.6
New Baghouse -
Capital Cost ($/kW) 149.2 to 639.2
Annualized Cost (mills/kWh) 10.2 to 75.7
Cost Per ton of S02 Removed ($/ton) 522.2 to 8,632.4
Appendix F contains tables of costs as a function of boiler size, coal
sulfur, capacity factor, and retrofit factor.
2-67
-------
References
1. Miller, M. J. SO- and NO Retrofit Technologies Handbook. CS-4277-SR,
Electric Power Research Institute, Palo Alto, California, 1985, 366 pp.
2. Policy Analysis Department. Steam Electric Plant Factors, 1987 Edition.
The National Coal Association, Washington, D.C., 1987.
3. Energy Ventures Analysis, Inc. Coal Markets and Utilities' Compliance
Decisions. Palo Alto, CA. EPRI Report P-5444, September 1987.
4. Tennican, M. L., R. E. Wayland, and D. M. Weinstein. Agenda of Critical
Issues: Coal Price and Availability. EA-3750, Electric Power Research
Institute, Palo Alto, California, 1984, 41 pp.
5. Palmisano, P. J. and B. A. Laseke, User's Manual for the Integrated Air
Pollution Control System Design and Cost Estimating Model, Version II,
Volume I, EPA-600/8-86-031a (NTIS PB87-127767), September 1986.
7. E. H. Pechan and Associates, Inc. and PEI Associates, Inc. Acid
Deposition Control Techniques for the New England States. Northeast
States for Coordinated Air Use Management, Boston, Massachusetts, 1986.
8. Wisconsin Public Service Commission. Utility Sulfur Dioxide Cleanup -
Cost and Capability. Madison, Wisconsin, 1986, 116 pp.
9. Bechtel National, Inc. Advanced Physical Fine Coal Cleaning,
DOE/PC/81205-T6-Vol. 1 (DE89003925), September 1988, p. 1-1.
10. Kilgroe, J. D. and J. Strauss. Coal Cleaning Options for SO* Emission
Reduction. Power Magazine Conference on Acid Rain, Washington, D.C.,
1985, 53 pp.
11. McCandless, L. C. and R. B. Shaver. Assessment of Coal Cleaning
Technology, First Annual Report. EPA-600/7-78-150 (NTIS PB287-091), July
1978.
12.. Onursal, A. B., J. Buroff, and J. Strauss. Evaluation of Conventional and
Advanced Coal Cleaning Techniques. EPA-600/7-86-017 (NTIS PB87-104535),
September 1986.
13. Hucko, R. E., H. B. Galal, and P. S. Jacobsen, Status of DOE Sponsored
Advanced Coal Cleaning Processes. DOE/PETC/TR-89/3 (DE89008772),
March 1989. pp. 1, 2.
14. Cavallaro, J. A., M. T. Johnston, and A. W. Deurbrouck. Sulfur Reduction
Potential of the Coals of the United States. EPA-600/2-76-091 (NTIS
PB252-965), April 1976, 323 pp.
15. Holt, E. C. An Engineering/Economic Analysis of Coal Preparation Plant
Operation and Cost. EPA-600/7-78-124:: (NTIS PB285-251), U. S. Department
of Energy and U. S. Environmental Protection Agency, Washington, D. C.,
1978, 296 pp.
2-68
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16. Thompson, R. E. and M. W. McElroy. Guidelines for Retrofit Low NO
Combustion Control. In: Proceedings: 1985 Symposium on Stationary
Combustion N0v Control, Vol. 1, EPA-600/9-86-021a (NTIS PB86-22
5042), July 1986, p. 27-1.
17. Douglas, J. Retrofit Strategies for NOW Control. EPRI Journal, 12(2):
26-31, 1987. x
18. Keeth, R. J., et at. Economic Evaluation of FGD Systems (Volume 1).
CS3342, Electric Power Research Institute, Palo Alto, California, 1983,
400 pp.
19. Melia, M. T., R. S. McKibben, F. M. Jones and J. L. Kelly. Trends in
Commercial Applications of FGD. In: Proceedings: Tenth Symposium on Flue
Gas Desulfurization, Atlanta, Georgia, November 1986, Vol. 1,
EPA-600/9-87-004a (NTIS PB87-166609), February 1987, p. 2-23.
20. Liegois, W. A. and D. A. Wicks. Gypsum By-product FGD System.
In: Ninth Symposium on Flue Gas Desulfurization, Cincinnati, Ohio,
June 1985, Volume 2, EPA-600/9-85-033b (NTIS PB 86-138658), December 1985,
p. 716. .
21. Mzyk, D. and J. Zmuda. By-product Gypsum Production at a 2300 MW Power
Plant. In: Ninth Symposium on Flue Gas Desulfurization, Cincinnati,
Ohio, June 1985, Volume 2, EPA-600/9-85-033b (NTIS PB 86-138658),
December 1985, p. 730.
22. Ellison, W., and L. M. Luckevich. FGD Waste: Long-Term Liability or
Short-Term Asset? Power, 128(6)-.79-83,1984.
23. Ottmers, D. M., Jr., et al. Evaluation of Regenerate Flue Gas
Desulfurization Processes. RP 535-1, Electric Power Research Institute,
Palo Alto, California, 1976.
24. Hudak, C. E. and J. M. Burke. Sulfur Oxides Control Technology Series:
Flue Gas Desulfurization, Well man-Lord, Summary Report.
EPA-625/8-79-001, U.S. Environmental Protection Agency, 1979.
25. Behrens, G. P., et al. The Evaluation and Status of Flue Gas
Desulfurization Systems. CS-3322, Electric Power Research Institute,
Palo Alto, California, 1984, 702 pp.
26. Ellison, W., et al. West Germany Meets Strict Emission Codes by
Advancing FGD. Power, 130(2):29-33, 1986.
27. Erdman, D. A. Mag-Ox Scrubbing Experience at the Coal-Fired Dickerson
Station of Potomac Electric Power Company. In: Proceedings:
Symposium on Flue Gas Desulfurization, Atlanta, November 1974,
Volume II, EPA-650/2-74-126b (NTIS PB 242573), December 1974, p. 729.
28. Shattuck, et al. Retrofit FGD Cost Estimating Guidelines. Prepared by
Steams-Catalytic Corporation for the Electric Power Research Institute
(EPRI), Palo Alto, California, EPRI Report CS-3696, October 1984.
2-69
-------
29. Palazzolo, M. A. and M. A. Baviello. Status of Dry SO, Control Systems:
Fall 1983, EPA-600/7-84-086 (NTIS PB 84-232503), U. Sf Environmental
Protection Agency, Research. Triangle Park, North Carolina, August 1984.
30. Blythe, G. M., et al. Field Evaluation of a Utility Spray Dryer System.
EPRI Report CS-3954, Electric Power Research Institute, Palo Alto,
California, 1985, 272 pp.
31. Robards, R., et al. Spray Dryer/ESP Testing for Utility Retrofit
Applications on High-Sulfur Coal. In: Proceedings of the American Power
Conference, TVA/OP/ED4T86/8, Chicago, Illinois, 1986.
32. Blythe, G. M., et al. Evaluation of a 2.5-MW Spray Dryer/Fabric Filter
SO- Removal System. EPRI Report CS-3953, Electric Power Research
Institute, Palo Alto, California, 1985, 312 pp.
33. Donnelly, J. R. and K. S. Felsvang. Spray Dryer Absorption Applications
for High Sulfur Coal. In: Proceedings of the Third Annual Pittsburgh Coal
Conference, Pittsburgh, Pennsylvania, 1986, 974 pp.
34. Yen, J. T., et al. Performance of a Spray Dryer/ESP Flue Gas Cleanup
System During Testing at the Pittsburgh Energy Technology Center. In:
Ninth Symposium on Flue Gas Desulfurization, Cincinnati, Ohio, June 1985,
Volume 2, EPA-600/9-85-033b (NTIS PB 86-138658), December 1985, p. 670.
35. Brown, C. A. et al. Results From the TVA 10-MW Spray Dryer/ESP
Evaluation. In: Proceedings: First Combined FGD and Dry SO, Control
Symposium, EPA-600/9-89-036a (NTIS PB 89-172159), March 198$, p. 3-1.
36. Donnelly, J. R., S. K. Hansen, M. T. Quach, and P. S. Farber. Industrial
Spray Dryer FGD Experience. In: Proceedings of the 1986 ASME Industrial
Power Conference, Pittsburgh, Pennsylvania, 1986, pp. 71-76.
2-70
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SECTION 3
NEAR-COMMERCIAL TECHNOLOGIES
INTRODUCTION
This section discusses six technologies which are in the process of being
demonstrated by U. S. utilities or which have been commercially applied to
power plants overseas. The first two technologies, integrated gasification
combined-cycle (Section 3.1) and fluidized bed combustion (Section 3.2), are
integrated processes that can increase the electric generating capacity of a
plant while minimizing SO- and NO emissions. Because of their design and high
capital cost, these two repowering technologies will be used by utilities that
have a need to increase or maintain base load electric generating capacity as
well as to reduce emissions. When compared with the cost of a new plant
(including emission control systems), addition of base load generating capacity
through repowering of existing capacity can be very economical.
Post-combustion NO control technology (Section 3.3), has been
^
commercially applied to boilers burning low-sulfur fuels in Europe and Japan,
but has little or no long-term operating experience with medium- and
high-sulfur coals. Because of specific process concerns, it is categorized in
this report as a near-commercial technology until its use with higher sulfur
fuels has been demonstrated.
The final three technologies, furnace and duct (also known as hot-side and
cool-side) sorbent injection (Sections 3.4 and 3.5) and reburning (Section 3.6)
have limited commercial application. These technologies have.limited or
planned large-scale demonstration experience in the U. S. and should be
commercially available within the next few years.
3-1
-------
3.1 INTEGRATED GASIFICATION COMBINED CYCLE
Description
Integrated gasification combined cycle (IGCC) is an alternative to
conventional coal-fired electric power generation with post-combustion emission
controls. Because of its overall design, emissions of sulfur and nitrogen
oxides and participates from IGCC facilities are projected to be significantly
lower than from existing technologies. Figure 3.1-1 is a generalized block
flow diagram of an IGCC facility designed to 1) convert coal (via partial
oxidation and gasification) into a fuel gas, 2) remove reduced sulfur species
(primarily hydrogen sulfide) from the fuel gas, 3) use the clean fuel gas to
produce electricity in a combined-cycle application, and 4) treat and/or
recover byproducts from the waste streams generated. Existing technology for
removal of sulfur species from fuel gas requires gas cooling followed by
scrubbing with any of several proprietary chemical processes. Cooling the gas
imposes both additional capital cost and a significant penalty on process
thermal efficiency. Advanced processes for removal of H^S from hot fuel gas
(1,000 to 1,200°F) using adsorption by metal oxides are under development with
funding from DOE. Development of hot gas cleanup technology will simplify IGCC
process complexity and reduce costs.
A number of coal gasification processes have been examined for use in
combined-cycle applications. The most successful demonstration of IGCC
technology in the U. S. is the Cool Water Coal Gasification Demonstration
Project. The Cool Water facility, located in Daggett, California, is based on
use of Texaco gasification technology with cold gas cleanup. The facility is
designed to gasify 1000 tons of coal per day and produce approximately 100 MM
of net electrical power. The facility underwent successful startup in 1984 to
become the first successful domestic operation combining coal gasification and
combined cycle equipment. The description of Cool Water presented below is
representative of the general concept of IGCC facilities.
3-2
-------
OJ
f
CO
Air Ik-
Air Separation
Oxygen
Coal
Coal
Preparation
Steam
Steam
•»
-p»
Saturated Steam
Gasification
Acid Gas
Removal
Condensate Ash to
to Treatment Disposal
I
Flue Gases
Sulfur
Recovery
i
Heat Recovery
Steam Generation
Hot Turbine
Exhaust Gas
Clean
Fuel Gas
Combustion
Turbines
Air
Sulfur
By - Product
Superheated
Steam
Steam
Turbine
Electric
Power
Figure 3.1-1. Generalized block flow diagram of combined
cycle coal gasification power generation
-------
The Texaco gasifier is a pressurized, downflow, entrained reactor in which
pulverized coal is fed as a coal/water slurry. High purity oxygen
(95+ percent) is also introduced into the top of the gasifier. The coal
undergoes partia> oxidation at high temperature and pressure to produce a raw
fuel gas composed primarily of hydrogen, carbon monoxide, and carbon dioxide.
The raw fuel gas also contains reduced sulfur and nitrogen species, water
vapor, entrained molten slag, and soot particles; no condensible hydrocarbons
are found in the gas.
The raw fuel gas 1s cooled in radiant and convective coolers, producing
saturated steam for use In the combined-cycle power generation system. The
fuel gas is initially treated in a particulate scrubber and a series of heat
exchangers and coolers. Cooled fuel gas is then sent to the sulfur removal
unit (the Selexol process is used at Cool Water) with a design sulfur removal
efficiency of 97 percent. Removed sulfur compounds are subsequently converted
Into elemental sulfur 1n a Claus sulfur recovery unit equipped with a SCOT tail
gas treating unit.
The clean fuel gas is then reheated and saturated with moisture before
being combusted in a gas turbine to produce electricity. Moisture addition is
used to reduce NO formation during combustion. Exhaust gas from the gas
turbine is sent through a heat recovery steam generator (HRS6) to produce
steam. This steam plus that from the fuel gas heat exchangers and coolers is
sent to the steam turbine to produce additional electricity. Cooled flue gas
from the HRSG is discharged through a stack to the atmosphere.
Auxiliary processes included in an IGCC facility include 1) coal storage
and preparation, 2) oxygen production (if pure or nearly pure oxygen is used as
the gasifier reactant) and compression, 3) wastewater treatment, and
4) other processes normally present in a coal-fired electric generating plant
(e.g., cooling towers and ash handling/disposal).
Applicability
A number of coal gasification technologies are commercially available, and
oil- and gas-fired combined-cycle power plants are currently in commercial use.(l)
3-4
-------
With the recent success of the Cool Water Demonstration Program and
continued research/demonstration activities, IGCC has become a rapidly
emerging alternative for electricity generation at new plants and for
repowering existing plants. Potential economic advantages include
high modularity which will allow IGCC plants to be constructed in relatively
small increments and ability to use lower priced (higher sulfur) coals.
Retrofit to an existing coal-fired power plant will require construction of
new gasification and combustion turbine facilities. However, depending on
site-specific factors, it may be possible to incorporate some existing
equipment into the IGCC plant (e.g., steam turbines, boiler feed-water and
cooling systems, coal handling, and electric generator and auxiliaries).
Performance
The electric power industry has a growing interest in IGCC technology
because of a number of environmental advantages versus conventional
coal-fired power plants, including 1) lower SO-, NO , and particulate
emissions, 2) lower water consumption, and 3) lower land requirements.
Sulfur dioxide emissions from an IGCC facility primarily arise from two
sources: combustion flue gas and tail gases from the sulfur recovery unit.
Very stringent control of sulfur emissions from an IGCC plant is attainable
even for gasification of high-sulfur coals.
The sulfur removal unit in an IGCC is designed to remove sulfur species
(primarily hydrogen sulfide) from the fuel gas prior to combustion, rather
than removal of SO- following combustion. Sulfur control efficiencies
greater than 95 percent are achievable with a properly designed sulfur
recovery unit. Sulfur species remaining in the fuel gas are converted to
sulfur oxides in the gas turbine combustor.
Sulfur emission test results from Cool Water are summarized in
Table 3.1-1. These results indicate that S02 emissions from IGCC can be
reduced to levels significantly below current utility NSPS requirements.
3-5
-------
TABLE 3.1-1. AVAILABLE SO EMISSIONS DATA FOR COOL WATER DEMONSTRATION PROGRAH
Combined -Cycle Combustion Flue Cases
Sulfur
Coal Gasified Removal (SO ppmvd") (Ib SO /hr> (Ib SO./ 10 Btu )
Lou sulfurC NR NR NR 0.032
Lou sulfur0 971 NR NR 0.034
Lou sulfur' NR fl. 59 33. l" NR
Incinerator Stack
(SO ppavd8) (Ib SO /hr) (Reference)
NR NR 2
NR NR It
54 3.2 7
NR = Not reported.
Parts per million by volume, dry basis.
Pounds SO per million Btu of coal.
Sulfur content not reported.
0.46 ppnvd of H,so, mist also reported.
2.7 Ib/hr of H SO mist also reported.
2 4
-------
Although the Cool Water test data shown are for gasification of a low-sulfur
coal, similar results are expected with medium- and high-sulfur coals.
The Electric Power Research Institute (EPRI) has funded a number of
studies relating to IGCC technologies (3,5,6,8,10). As part of these
studies, estimates of the performance and cost of IGCC facilities gasifying
high-sulfur Illinois #6 coal have been developed. Table 3.1-2 summarizes
the sulfur control levels incorporated into these designs. Even higher
levels of sulfur removal (97-99 percent) are achievable at small increases
in overall capital cost.
Nitrogen oxide emissions from IGCC facilities are primarily associated
with the gas turbine combustion gases. Test data from Cool Water,
summarized in Table 3.1-3, indicate NO emissions equal to one-tenth the
current utility NSPS for coal-fired boilers are achievable. As discussed
previously, the Cool Water facility injects steam into the fuel gas prior to
combustion in the gas turbine to reduce NO emissions.
Cost
Cost and emission data for several IGCC power plants are summarized in
Table 3.1-4. Additional cost data compiled by EPRI (9) provide capital cost
estimates of $1,204-1,576 per kW (in December 1984 $) for a 500 MW plant
using bituminous coal. Key cost variables were the gasification technology
(Texaco, Shell, KRW, and British Gas/Lurgi), method of gas cooling (none of
these include hot-gas cleanup), and turbine design (current or advanced).
The following table presents the range of IAPCS values for capital and
annualized costs. The lower capital and annual1zed costs are for a large
unit with a high capacity factor and high heat rate, while the higher
capital and annualized costs are for a small unit with a low capacity factor
and low heat rate. Figures 3.1-2 and 3.1-3 show these costs as a function
of boiler size and heat rate (cost bases differ from the costs presented in
Table 3.1-4). Reported heat rate values (9) range from 7,130 Btu/kWh for
Texaco fuel cell combined cycle to 10,250 Btu/kWh for Texaco gasification
combined cycle burning bituminous coal. For this study two cases were
considered for cost comparison, 8,000 and 10,000 Btu/kWh heat
3-7
-------
TABLE 3.1-2. DESIGN SULFUR EMISSIONS INFORMATION FOR EPRI IGCC STUDIES
Gasification
Technology
Texaco
British Gas Corp/
Lurgi (slagging)
Shelld
Kellogg Rust f
Westinghouse
Feed
Type
111. #6
111. #6
111. #6
111. #6
Coal
Sulfur
Content
3.4%
3.4%
3.4%
3.4%
Overall
Sulfur Removed
95%
94.4%
90%
95%
Sulfur Dioxide
Emissions, lb/10 Btua
0.28
0.34
0.56
0.28
aBtu of coal feed.
Reference 3.
d
cReference 5.
Reference 6.
ePlant was designed to meet NSPS (i.e., 90% control).
Reference 8.
3-8
-------
TABLE 3.1-3. AVAILABLE NOV EMISSIONS DATA FOR COOL WATER
DEMONSTRATION PROGRAM
Reference
2
4
7d
NR = Not reported.
aPer million Btu of
Parts per million
and 15% 02.
lb/106 Btua
0.058
0.059
NR
coal .
by volume, dry basis;
ppmvd
23
NR
22.8/20.8
normalized to
lb/hrc
61
NR
61.2/57.2
ISO humidity (60%)
Calculated as N02.
Results shown are for two test periods.
3-9
-------
TABLE 3.1-4. SUMMARY OF REFERENCE INFORMATION OH EMISSION AND COST DATA FOR IGCC FACILITIES
Plant SO SO Emissions
Facility Size, MU Removal Ib/MBtu
Texaco (3)
Radiant 600 95X 0.2B
Radiant & Corwectlve 589 95X 0.28
Total Quench 571 95X 0.28
Shell (a) 1.122 90X 0.56
Kellogg Rust 547 95X 0.28
Uestlnghouse (8)
Capital Annual
NO Emissions Requirement O&M Cost. Cost
X
Ib/NBtu Millions S/kV SI. 000 Basis
0.04 . 820 1,360 24,800 1/83 S
01.3 860 1,460 26,100 1/83 S
0.04 740 1,300 22,700 1/83 S
NR 1,420 1,260 43,300 1981 S
<75C 805 1,470 26.800 1/83 S
Reflects cost of entire plant.
Based on 100X capacity and combination of Illinois 06 coal.
Units « ppnvd, parts per million volume (dry basis).
-------
Heat Rate = 8,000 Btu/kWh
Heat Rote =10,000 Btu/kWh
Capacity Factor = 50%
100
1.300
C0
Figure 3.1-2. IGCC - Capital Cost, Current 1988 Dollars
210
Heot Rate = 8.000 Btu/kWh
Heat Rate = 10,000 Btu/kWh
Capacity Factor = 50%
100
300
500
700
900
1,100
1,300
MW
Figure 3.1-3. IGCC - Levelized Annual Cost,
Current 1988 Dollars
3-11
-------
rate, with direct capital costs for both cases assumed to be the same. This
is not to be confused with the more advanced gasification process but is
simply to present costs for a range of heat rates.reported. The total
capital cost for the low heat rate value is considerably lower due to the
lower preproduction and inventory costs.
Range
Capital Cost ($/kW) 1,344 to 2,919
Annualized Cost (mills/kWh) 80.3 to 206.0
Appendix 0 contains tables of costs as a function of boiler size, capacity
factor, and boiler heat rate.
Because IGCC is an integrated process producing electricity as well as
reducing SO, and NO emissions, it is difficult to separate costs associated
£ A
with emission reduction processes. However, for one facility the capital
.cost for acid gas removal and sulfur recovery were calculated to be
approximately 6 percent of the total plant capital investment.. Additional
cost information may be found in references 3, 6, and 8.
3-12
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3.2 FLUIDIZED BED COMBUSTION
Description
Fluidized bed combustion (FBC) is an integrated technology for reducing
SOg and NOX emissions during combustion of coal. As with IGCC, FBC can be
installed at new plants or used to repower existing plants. FBC
technologies based on operation at atmospheric and pressurized conditions
have been developed. Atmospheric FBC (AFBC) systems are similar to a
conventional boiler in that the furnace operates at near atmospheric
pressure and depends upon heat transfer of a working fluid (i.e., water) to
recover the heat of combustion. Pressurized FBC operates at pressures
greater than atmospheric pressure and recovers energy through both heat
transfer to a working fluid and use of the pressurized gas to power a gas
turbine. The pressurized systems offer the potential for smaller equipment
sizes (and thus simpler retrofits), lower capital cost, and higher
efficiency.
Atmospheric FBC --
Figure 3.2-1 presents a simplified AFBC process flow diagram. Coal and
limestone are fed into a bed of hot particles (1400-1600°F) fluidized by
upflowing air. SOg formed during combustion reacts with the calcined lime-
stone to form calcium sulfate, thus reducing SCL emissions and avoiding the
need for post-combustion controls. The relatively low combustion
temperature limits NO formation, reduces ash fusion problems, and optimizes
sulfur capture. One of the major advantages of AFBC technology is its
ability to burn a wide range of fuels, including those with low heat and
%
high sulfur contents.
There are two major AFBC types: the bubbling or dense bed, and the
circulating or dilute bed. In the bubbling bed system, coal and limestone
are continuously fed into the boiler from over or under the bed. The bed
materials consisting of unreacted, calcined, and sulfated limestone, coal,
and ash are suspended by the combustion air blowing upwards through the air
distributor plate. Bed material is drained from the bed to maintain the
3-13
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Convection
i
Pass
Coal Limestone
Freeboard «
Splash Zone «
Bed «
Transport Air
Fluldlzlng Air
Forced Draft Air
Flue Gas
Cyclone
Recycle
Distributor
Plate
Plenum
Waste
Waste
| Compressor |
Figure 3.2-1. Simplified AFBC process flow diagram.
3-14
-------
desired bed depth. Some bed material is entrained in the upflowing flue gas
and escapes the combustion chamber. Approximately 80 to 90 percent of this
fly ash is collected in the cyclone and is then either discarded or
reinjected into the bed. Reinjection of ash is useful in improving
combustion efficiency and limestone utilization. In general, combustion
efficiency increases with longer freeboard residence times and greater ash
recycle rates. Fly ash not collected in the cyclone is removed from the
flue gas by an ESP or fabric filter.
Circulating fluidized bed is a more recent development in AFBC
technology. The two major differences between circulating and bubbling
AFBC's are: the size of the limestone particles fed to the system, and the
velocity of the fluidizing air stream. Limestone feed to a bubbling bed is
generally less than 0.1 inches in size, whereas circulating beds use much
finer limestone particles, generally less than 0.01 inches. The bubbling
bed also incorporates relatively low superficial air velocities, ranging
from 4 to 12 ft/sec. This creates a relatively stable fluidized bed of
solid particles with a well-defined upper surface. Circulating beds have
superficial velocities ranging from 20 to 40 ft/sec. As a result, a
physically well-defined bed is not formed; instead, solid particles (coal,
limestone, ash, sulfated limestone, etc.) are entrained in the transport
air/combustion gases. These solids are then separated from the combustion
gases by a cyclone or other separating device and circulated back into the
combustion region, along with fresh coal and limestone. A portion of the
collected solids are continuously removed from the system to maintain
material balances. Circulating beds are characterized by very high
recirculated solids flow rates, up to three orders of magnitude higher than
the combined coal/limestone feed rate (11).
Several advantages of circulating bed over bubbling bed design have
been identified by FBC unit vendors (due to limited test and commercial
operating data, however, not all of the advantages have been fully
demonstrated):
t Higher combustion efficiency, exceeding 99 percent;
• Greater limestone utilization, due to high recycle of unreacted
sorbent and small limestone feed size (greater than 85 percent SO-
3-15
-------
removal efficiency is projected with a Ca/S ratio of about 1.5,
with the potential for greater than 95 percent SO- removal
efficiency);
t Lower NO emissions because of staged combustion (less than
100 ppm NO are projected);
A
t Potentially fewer corrosion and erosion problems, since the heat
transfer surface is less likely to be located in the primary
combustion zone;
• Minimal excess air requirements, since the high velocities promote
good mixing and combustion efficiency;
• Less dependence on limestone type, since reactivity is improved
with the fine particle sizes; and
• Reduced solid waste generation rates, because of lower limestone
requirements.
Potential disadvantages are:
• Increased capital costs;
• Greater energy losses due to high pressure drops;
• Increased combustor height compared to a bubbling bed;
t Uncertainty regarding the hot cyclone's ability to adequately
separate solids and gases and to resist erosion and corrosion;
• Difficulty in scaling up for large units (over approximately
200 MW), limiting application generally to smaller boiler sizes;
• Less ability to use existing boiler equipment in retrofit
applications; and
• Erosion of components subjected to impingement of high
velocity particles.
Table 3.2-1 presents the estimated operating data for a 200 MW AFBC.
Table 3.2-2 summarizes full-scale utility AFBC demonstration facilities
3-16
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planned for startup between 1986 and the early 1990's. During retrofit of
Northern States Power's Black Dog Unit No. 2, generating capacity was
increased from 85 MW up to 125 MW by increasing operating pressures and
modifying the turbine-generator as part of the overall retrofit program.
Estimated cost was $58 million ($460/kW). Retrofitting AFBC on the four
units at Wisconsin Electric's Oak Creek Station is expected to cost an
estimated $380 million ($760/kW). The projected output includes recovering
130 MW of generating capacity lost due to coal changes and mechanical
limitations on the existing boilers (12). The recovery of generating
capacity lost during previous plant derates or expansion in capacity due to
overall plant upgrading is one of the key economic benefits from repowering
an existing plant with FBC.
TABLE 3.2-1. ESTIMATED OPERATING DATA FOR 200 MW AFBC
Performance at Full Load
Steam Cycle Heat Rate, Btu/kWh 7724
Boiler Efficiency, percent 87.5
Gross Plant Heat Rate, Btu/kWh 8823
Auxiliary Power, MW 14.9
Net Plant Heat Rate, Btu/kWh 9500
Turbine Generator Gross Output, MW 209.6
Turbine Generator Net Output, MW 194.7
Consumption/Production @ 65% CF
Coal, 103 TPY 521
Limestone 9 Ca/S = 2.5, 103 TPY 173
Waste Solid, 103 TPY 229
Source: Reference 13.
3-17
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TABLE 3.2-2. FULL-SCALE UTILITY AFBC DEMONSTRATIONS
Utility
Colorado Ute
Montana-Dakota
Station
Nucla
Heskett
New/
Retrofit
New
Retrofit
Bed Type
Circulating
Bubbling
Size
(MW)
110
85
Start-Up
1987
1987
Utilities
Northern States
Power
Tennessee Valley
Black Dog
Shawnee
Retrofit
New
Bubbling
Bubbling
125
160
1986
1988
Pressurized FBC --
Pressurized fluidized bed combustion (PFBC) is similar to AFBC with the
exception that combustion occurs under pressure. By operating at pressure, it
is possible to reduce the size of the combustion chamber and to develop a
combined-cycle or turbocharged boiler capable of operation at higher efficien-
cies than atmospheric systems. The turbocharged boiler approach recovers most
of the heat from the boiler through a conventional steam cycle, leaving only
sufficient energy in the gas to drive a gas turbine to pressurize the
combustion air. .The combined cycle system extracts most of the system's energy
through a gas turbine followed by a heat recovery steam generator and steam
turbine. As with IGCC (Section 3.1), hot gas cleanup technology is a critical
need in combined cycle PFBC development.
The smaller size of a PFBC boiler may be especially attractive in retrofit
applications where space is limited. However, PFBC 1s not as well developed as
3-18
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AFBC. To date, work on PFBC has been limited mainly to component and
pilot-scale testing at facilities in the U. S., England, and Sweden.
Commercial PFBC systems have been ordered by two European utilities. There is
currently one utility boiler size PFBC unit (75 MW) under construction in the
United States (Tidd) (14) with another one planned for 1995 (Philip Sporn).
Initial utility PFBC demonstration projects are not expected to be in operation
until the early 1990's with earliest commerical availability projected in the
1995-2000 time frame (15).
Applicability
For new plants, AFBC can be applied to nearly any plant site because of
the inherently wide design ranges for fuel and operating conditions. Single
unit AFBC design is presently limited in size to about 300 MW. At these sizes,
unit design can be modularized to reduce construction cost. Construction times
of 2-3 years are possible, thus allowing the utility to schedule construction
in response to load growth or other requirements.
Retrofitting AFBC at an existing plant requires either 1) installation of
a totally new boiler to supply steam to an existing turbine or 2) replacement
of the lower furnace section of the boiler while continuing to use portions of
the convective heat transfer sections of the existing boiler. Major design
questions in making this decision are the layout, condition, and operating
limits (e.g., pressure/temperature) of the existing boiler and particulate
control system and the availability of additional space if an entire new boiler
is installed.
Current AFBC designs are limited to heat release rates of
2
1.0 million Btu/ft . Therefore, an existing boiler with a heat release rate of
2
1.5 million Btu/ft would either require an AFBC approximately 50 percent
larger than the existing boiler plan area or have to be derated to two-thirds
of its present capacity (16). Based on this, retrofit of AFBC will likely be
2
limited to boilers with heat release rates below 1.5 million Btu/ft . Because
3-19
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most cyclone-fired and wet-bottom pulverized coal boilers have higher heat
releases, these units are not likely candidates for AFBC retrofit. Bubbling
beds are preferred for retrofit applications because they require less space
and can use more of the existing plant equipment.
Performance
Extensive testing of AFBC has been performed in numerous vendor pilot
plants, full-scale Industrial AFBC's, the EPRI/Babcock & Wilcox test facility
in Alliance, Ohio (rated at 20,000 Ib/hr of steam), the 20 MW Tennessee Valley
Authority pilot plant at Shawnee Station, and the 15 MW Northern States Power
French Island Pilot Plant. The EPRI/B&U facility has unique scaling features
for commercial units, such as a high free board zone, to simulate utility
boiler residence times and temperatures. Test results have shown overall
combustion efficiency of 98 percent and sulfur capture of 90 percent with a
calcium-to-sulfur ratio of 2 to 1. NOV emission levels achievable with AFBC
6
are estimated at less than 0.4 lb/10 Btu.
The 20 MW TVA pilot plant has been used to evaluate a number of areas
requiring technological developments related to size scale-up and long-term
operation of commercial AFBC systems. Objectives of the 20 MW pilot test,
cosponsored by TVA and EPRI, included:
• Demonstrate acceptable, long-term operation and performance
(efficiency, reliability, emissions control, etc.);
t Develop and demonstrate automatic control and load following
capability;
• Develop design correlations for confident scaling to commercial size
boilers;
0 Demonstrate the environmental control capability of the unit as a
basis for assessing the environmental acceptability of AFBC on a
commercial scale;
0 Develop the capability to specify and design critical auxiliary
equipment, especially the coal, limestone, and recycle feed systems;
3-20
-------
• Test and evaluate materials of construction for the boiler and its
auxiliaries; and
• Train engineers and operators for future units.
Performance results were favorable. Combustion efficiency, sulfur
capture, and NO emissions have been well within expectations. The coal feed
systems were the largest problem. The underbed feed system demonstrated
high combustion and sulfur control efficiencies, but required a more complex
feed system design and operated less reliably. The overbed feed system
demonstrated high equipment reliability but less acceptable process
efficiencies, especially when the coal contained significant (15-20 percent)
fines. Remedies for these problems were investigated prior to the final design
of the 160 MW TVA demonstration project.
Flyash recycle is another key design factor for sulfur capture and
combustion performance optimization. As previously mentioned, combustion
efficiency and limestone utilization increase with greater recycle. Other
methods for improving limestone utilization include reinjection of sulfated
limestone after pulverizing and hydration.
Cost
Analysis of system economies for new plants have generally found FBC to be
cost competitive with pulverized coal firing combined with limestone FGD. Cost
estimates by EPRI for a conventional boiler, AFBC, and PFBC are presented in
Table 3.2-3 for a new plant.
The following table presents the range of values estimated using IAPCS for
capital and annualized costs of new AFBC plants. The lower capital cost is for
a large unit with a high capacity factor, while the higher capital cost is for
a small unit with a low capacity factor. Figures 3.2-2 and 3.2-3 show these
costs as a function of boiler size and heat rate. The reason for lower costs
in the lower heat rate value cases is due to the lower preproduction and
inventory costs.
Range
Capital Cost ($/kW) 1,132 to 2,382
Annualized Cost (mills/kWh) 77.0 to 182.4
3-21
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TABLE 3.2-3. COMPARATIVE ECONOMICS OF NEW 500 MW CONVENTIONAL AND
FBC POWER PLANTS3 (1988 dollars)
Conventional
Boiler w/ . Turbocharged Combined-Cycle
LS FGD AFBC PFBCD PFBC
Capital Cost, S/kW 1,233 1,184 1,247 1,394
Heat Rate, Btu/kWh 10,060 10,000 . 9,703 8,980
Total Cost, Mills/kWh
Constant $ 37.3 23.0 36.9 39.4
Current $ 64.2 62.4 63.5 67.8
aBased on a 500 MW boiler with 3.5 percent sulfur coal at $1.55/million Btu and
capacity factor of 65 percent.
bScaled from 250 MW based on a scaling factor of 0.8; circulating bed
design.
Source: Reference 9.
3-22
-------
03
0.0
Heot Rate = 9,000 Btu/kWh
Heat Rate = 11,000 Btu/kWh
Capacity Factor = 50%
100
1.300
Figure 3.2-2. AFBC - Capital Cost, Current 1988 Dollars
x
*
ec
UJ
Q.
Heot Rote = 9,000 Btu/kWh
Heot Rote = 11.000 Btu/kWh
Copocity Factor = 50%
'90
100
I I T T
900 1.100
1.300
Figure 3.2-3. AFBC - Levelized Annual Cost, Current 1988 Dollars
3-23
-------
Appendix E contains tables of costs as a function of boiler size,
capacity factor, and boiler heat rate. Detailed costs are not available at
this time for PFBC, and therefore these costs are not presented.
3-24
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3.3 POST-COMBUSTION NOx CONTROL
Description
Selective Catalytic Reduction--
Selective catalytic reduction (SCR) is a flue gas treatment process for
removal of nitrogen oxides (NOX). In this process, gaseous ammonia diluted
with either air or steam is injected into the flue gas upstream of the SCR
catalyst. The ammonia/flue gas mixture then enters the catalyst where the
ammonia reduces NO to elemental nitrogen and water (17).
4NO + 4NH3 + 02 > 4N2 + 6H20
2N02 + 4NH3 + 02 > 3N2 + 6H20
Because the NO composition of the flue gas from combustion sources is
primarily NO, the stoichiometric NH,/NO mole ratio is approximately 1:1.
A
An operating temperature of 570-750 F is required to catalyze NO reduction.
^
Below this temperature range, the reaction between ammonia and NO slows
significantly. Above this temperature range, ammonia is oxidized to NO and
the catalyst can be damaged thermally.
There are two possible SCR configurations. The first design is a
hot-side or high-dust SCR where the SCR system is located between the
economizer and air preheater. The second design is a cold-side or low-dust
SCR where the SCR is located typically downstream of the particulate control
device and possibly downstream of the FGD system. Figure 3.3-1 presents
both SCR configurations. An economizer bypass can be incorporated into the
system to control low catalyst temperature excursions.
SCR catalyst formulations are typically based on oxides of titanium and
vanadium. Other active metals such as platinum, palladium, and copper can
also be used. In addition to reducing NO to N?, the catalyst also promotes
A £
the oxidation of SO- to S03 which in turn reacts with NH3 and H20 to form
ammonium bisulfate (NH^HSO^) as the flue gas cools. Ammonium bisulfate
deactivates the SCR catalyst by coating active catalyst sites and the air
3-25
-------
Cold Sioo S«>tem
Ash
Hot Siao System
Figure 3.3-1. Possible SCR Configurations.
3-26
-------
preheater elements reducing heat transfer efficiency and increasing metal
corrosion.
Catalyst elements can be produced in two general forms: composed
solely of catalyst material (homogeneous) or composed of either a ceramic or
metallic substrate that is coated with catalyst material (heterogeneous).
Each type of catalyst element has distinct advantages. For example, the
homogeneous element will not lose activity even if the outer layer is lost
due to erosion by flyash; however, eventually the physical strength can be
affected. To overcome catalyst degradation, the inlet face of the catalyst
can be protected with a hardened inert material. Coated catalyst elements
also have the potential for erosion.
The coated element will not be structurally weakened, but catalytic
activity can be lost. Metallic substrates are more resistant to breakage
and cause less pressure drop than a ceramic substrate due to thinner walls.
Alternatively, the ceramic material has better catalyst bonding properties,
is acid resistant, and lightweight.
Honeycomb, plate, and pipe catalyst supports are preferred for use with
coal-fired applications. A number of tests were performed on honeycomb and
plate catalysts (18). The honeycomb shape is the most common for reasons of
strength and ease of handling. The honeycomb shape consists of a square
block with parallel channels passing through it. The block is typically
6-20 inches on a side and up to 40 inches in length. The channels can be
either square, hexagonal, or triangular in shape. Plugging by particulates
is typically not a major problem when a vertical downflow reactor
arrangement and regular soot blowing are used.
Selective Non-Catalytic Reduction--
Selective non-catalytic reduction (SNR) is an alternate approach for
post-combustion reduction of NO to N-. The primary SNR technologies
involve injection of ammonia or urea (decomposes to ammonia) into the
convective section of the boiler at 1600-2000°F; no catalyst is used. At
temperatures above and below this range, ammonia is oxidized to NOX or exits
as unreacted NH,, respectively. In general, because of the short residence
time at these temperatures in the convective section and the difficulty of
thorough ammonia/flue gas mixing, NO reductions are lower (40-80 percent)
\.
£
3-27
-------
and ammonia emissions are higher than with SCR. Because a catalyst Is not
required, SNR costs are significantly lower than for SCR. However, because
of the lower performance level and limited experience with SNR on coal-fired
boilers, the rest of this section is limited to SCR technology.
Applicability
Since the SCR process requires flue gas at a specific temperature, it
is simpler to apply it to new boilers than to existing boilers. In new
boilers, the SCR system is normally placed between the economizer and air
preheater as economizer exit temperatures of utility boilers are typically
in the range of optimum SCR operation. Retrofit installations will require
flow modifications and additional ductwork to divert and return flue gas to
the existing flue gas handling system. Alternate arrangements include
installation of the reactor vessel downstream of a hot-side participate
control system or, by combining with flue gas reheat, downstream of the
existing emission controls for SO. and participates. This last approach may
be the most suitable for many retrofits due to boiler limitations.
SCR is considered a commercial technology; however, it has not been
widely applied in the U. S. More than 30,000 megawatts of coal-fired
utility boiler SCR systems have been operated in Japan and West Germany.
Three coal-fired pilot SCR demonstration projects have been conducted in the
U. S. The only full-scale U.S. application to a utility boiler is a gas-
and oil-fired demonstration unit at Southern California Edison's Huntingdon
Station. There are also three industrial boilers and a number of gas
turbines and process heaters using SCR for NO control, all in California.
Host of these applications involve very low-sulfur fuels; this
applicability of SCR technology to high-sulfur fuels has not been
commercially demonstrated to date. A primary concern is the impact of high
SO* levels on catalyst and air preheater plugging and corrosion due to
formation of ammonium bisulfate. Air preheater plugging is most severe in
units which fire high-sulfur fuels and remove particulates upstream of the
NO reactor. When particulates are removed downstream, plugging problems
are significantly reduced. It is believed that the particulates produce a
sandblasting effect that cleans the air preheater elements and that some of
3-28
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the NH4HS04 condenses on the flyash particles rather than the air preheater.
Air preheater plugging can also be reduced by limiting NH- slip to about 3
to 5 ppm (NH3 slip is the reacted NH, exiting the SCR reactor) and
redesigning the air heater. However, reducing NH3 slip will require
increasing catalyst volume if a specific NOX removal rate is to be
maintained. Off-line water washing of the air preheater may also be
required to minimize pressure drop; soot blowers are generally ineffective
at removing NH^HSO^ deposits from the air heater on-line.
NH. from an SCR reactor does not impair FGD SO, removal performance.
However, in some cases, the wastewater may require treatment to remove
nitrogen (nitrate) compounds. An unresolved concern is the potential for
NH, slippage or NH.HSO. to affect the performance of a downstream
particulate control device (e.g., by increasing pressure drop across a
downstream fabric filter or ESP).
Performance
SCR is capable of 80 to 90 percent NO reduction. The primary
variables which determine the amount of NO removed are the amounts of
^
ammonia and catalyst used; increasing ammonia and catalyst yields higher
removals. However, increasing the amount of ammonia injected also increases
ammonia slip. For this reason, the NH3 slip limit set for the initial NOX
concentration is the key catalyst sizing parameter, rather than the required
NOX reduction.
Catalyst bed design is usually specified in terms of a space velocity
or area velocity. The space velocity is the flue gas volumetric flow rate
divided by the catalyst bulk volume, while area velocity is flue gas
volumetric flow rate divided by the active catalyst area. High NO
reductions require low space (or area) velocities and, hence, large
quantities of catalyst. Low velocities are also required to control the
amount of NH, slip. To reduce NH, slip to 3-5 ppm levels, space velocities
o i j
in the range of 2,000-2,500 hr are generally required. Furthermore, NH3
slip increases dramatically at NH3:NOX injection ratios over 0.8 for high
space velocities and 0.9 for low space velocities. This constraint
contributes in holding SCR to a practical NO removal limit of about 80
3-29
-------
percent. The exact relationships between velocity and NOV reduction and NH,
A . O
emissions, however, are specific to the proprietary catalysts used by the
various SCR manufacturers.
Cost
The capital cost of an SCR system is a function of several factors
including NO reduction level, flue gas flow rate, fuel type and sulfur
content, and retrofit/new installation (19). Operating costs are primarily
a function of catalyst life.
The following table presents the range of values for capital cost,
annualized cost, and cost per ton of pollutant removed for hot side SCR
systems using IAPCS. The lower costs are for a large unit with a low
retrofit factor and the higher costs are for a small unit with a high
retrofit factor. Figures 3.3-2 through 3.3-4 show these costs as a function
of boiler size and catalyst life.
Ranoe
Capital Cost ($/kW) 85.4 to 240.6
Annualized Cost (mills/kWh) 6.1 to 39.4
Cost Per ton of NOX Removed ($/ton) 1,856 to 12,334.1
Estimates of catalyst life vary from 1-5 years. For the purpose of this
analysis, 1 and 3 year catalyst life is assumed. Appendix G contains tables of
costs as a function of boiler size, retrofit difficulty, capacity factor, and
catalyst life.
3-30
-------
o
Q.
z
JC
oc
Ul
Q.
150
140
130 -i
120 -
Legend
Catalyst Life = 1 yr or 3 yrs
Capacity Factor = 50%
Catalyst Cost = $23.342/ton
110 -
100 -
100
300
1.100
1.300
Figure 3.3-2. Selective Catalytic Reduction - Capital Cost
Current 1988 Dollars
19
18 -
17 -
16 -
15 -
14 -
13 -
12—
Legend
• Catalyst Life = lyr
+ Catalyst Life = 3yr
Capacity Factor = 50/5
Catalyst Cost = $23,342/ton
100
300
500
MW
700
900
1.100
1.300
Figure 3.3-3. Selective Catalytic Reduction - Levelized
Annual Cost, Current 1988 Dollars
3-31
-------
Q.O
5.4
5.2 -
5 -
4.8 -
4.6 -
4.4 -
4.2 -
4 -
3.8 -
3.6 -
3.4
3.2 -
3 -
2.8 -
2.6 -
2.4 -
100
Legend
• Cotolyst Life = 1yr
•*• Catalyst Life = 3yr
Capacity Factor = 50% .
Cotolyst Cost = S23,342/ton
300
500
700
MW
900
1.100
1.300
Figure 3.3-4. Selective Catalytic Reduction - Cost Per Ton
of NOx Removed, Current 1988 Dollars
3-32
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3.4 FURNACE SORBENT INJECTION
Description
Furnace sorbent injection (FSI, or LIMB) reduces S02 emissions by
injecting powdered limestone or hydrated lime into the upper furnace of a
coal-fired boiler. These sorbents decompose at high temperature to form
calcium oxide which reacts with SO. to form calcium sulfate (CaS03) (20).
Limestones ranging from calcitic (mainly CaC03, very low in MgC03) to dolomitic
(containing roughly equal amounts of Ca and Mg) have been tested. Atmospheric
and pressure-hydrated limes have also been used. The extent of S02 removal
that can be achieved depends on the flue gas composition, temperature, and
quench rate at the point of sorbent injection; sorbent composition and surface
area; and the calcium-to-sulfur ratio (Ca/S) and degree of mixing between the
sorbent and S02 in the flue gas (21).
The resulting calcium sulfate and unreacted sorbent are collected with the
flyash in a baghouse or electrostatic precipitator (ESP). Collected sorbent
and flyash can be either recycled to the furnace to increase sorbent
utilization or discharged to the solid waste disposal facilities. Figure 3.4-1
is a generalized diagram of the FSI process.
The advanced silicate (ADVACATE) process (22) can be added to FSI (or
stand alone process) for better SO- removal efficiency. Preliminary pilot
tests on ADVACATE system have resulted in 90+ percent S02 removal with a
calcium to sulfur ratio near 1.0. In this process cyclones may be used to
knock out coarse dust upstream of the ESPs. Part of the high calcium flyash in
the front section of the ESPs, silicate, is slurried under temperature and
pressure. Silica in the fly ash dissolves and reacts with calcium to form high
surface area silicates. The slurry is mixed with the remaining high calcium
fly ash to form a damp powder, which is then injected into the ductwork.
Performance
The most important variable affecting the calcination and sulfation
reactions is the gas temperature at the point of sorbent injection. As shown
in Figure 3.4-2, a peak in S02 removal occurs for most sorbents in the range
3-33
-------
Sorbent
Figure 3.4-1. Simplified schematic of furnace sorbent injection.
100
35
I
2
8"
80 -
60 -
40 -
20 -
Genstar Pressure
Hydrated Oolomitlc Lima
(1)
Longview Lime
(2)
Marianna
Limestone
(2)
St. Genevieve
Limestone
(2)
Ca/S in 0
1400
1800
2200
2600
Maximum Injection Temperature, F
Figure 3.4-2. Peak sorbent reactivity as a function
of temperature (2 3).
3-34
-------
between 1800°F and 2200°F. In most boilers, these temperatures are found in
the upper furnace. Because limited residence time is available in this zone to
allow thorough mixing of the sorbent and combustion gases, rapid distribution
of the sorbent over the full boiler cross-section is essential. At
temperatures higher than 2200°F, sorbent deactivation (due to sintering of the
sorbent surface and instability of the reaction products) lowers S02 removal.
At temperatures below 1800°F, reaction rates for both calcination and sulfation
are significantly reduced.
The sorbent quench rate (i.e., the rate of change in the flue gas
temperature after the point of sorbent injection) also affects SO- removal,
with SO. removal decreasing as quench rate increases (24). As quench rates
increase, sorbent residence time in the 1800-2200°F window is correspondingly
reduced. Expected SO. capture as a function of quench rate, Ca/S ratio, and
sorbent are shown in Table 3.4-1. This factor is especially significant for
cyclone-fired boilers for which quench rates in the upper furnace can approach
1500°F/sec.
Sorbent characteristics such as chemical composition and surface area also
affect the sorbent injection rate needed to achieve good SO. removal. The
surface area of several sorbents evaluated in a pilot-scale study are shown in
Figure 3.4-3. Pressure-hydrated dolomitic lime was the most effective for
removing S02; limestone was the least effective. Techniques that have been
studied for improving sorbent surface area include thermal pretreatment and use
of chemical promoters. Although these techniques have increased measured
surface area, pilot-scale combustion tests have not conclusively demonstrated
their effectiveness in improving SO- removal.
Sorbent residence time and distribution in the flue gas also impact S02
removal. As shown in Figure 3.4-4, longer residence times give the sorbent and
SO. more time to react and increase S02 removal. However, for retrofit
applications, the residence time of the boiler is already established and
cannot be readily varied. Optimal location, number, and type of injection
nozzles can maximize mixing of the flue gas and sorbent.
3-35
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TABLE 3.4-1. SO, CAPTURE AS A FUNCTION OF Ca/S RATIO, QUENCH RATE, AND
' SORBENT
Ca/S
Quench rate = 900
limestone 15-16 26-29 35-43 42-56
hydrate 19-20 35-40 53-56 69-70
CPHC 27-33 48-54 68-71 88-94
Quench rate = 700
limestone 17-19 29-31 38-44 45-57
hydrate 24-26 40-46 57-62 73-76
CPH 33-38 54-62 75-79 92^95
Quench rate =500
limestone 18-22 31-33 41-45 48-58
hydrate 28-32 45-52 61-68 77-82
CPH 40-43 60-70 81-89 95+
Quench rate = 300
limestone 19-25 32-36 44-46 51-59
hydrate 33-38 49-58 66-74 82-88
CPH 46-48 67-78 88-95 95+
aSO- capture is expressed as a percentage.
Quench rate is expressed as °F per second for the sulfation "window" of 2200
to 1600°F.
CCPH = caldtic pressure hydrate.
°F
3-36
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o
CO
100
80 -
60 -
40
20 -
Pressure-Hydrated
Calcium Dolomitic Lime
Hydroxide
Marianna
Limestone
St. Genevieve
Limestone
10 20 30
Calcine surface area, m2/g
40
FIGURE 3.4-3. S02 removal as a function of calcine surface area (23).
Residence Time
1.54 sec
0123456
FIGURE 3.4-4. S02 removal as a function of Ca/S ratio and residence time
using limestone (23).
3-37
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Three commercial-scale utility projects in the U. S. are demonstrating
furnace sorbent injection on eastern bituminous coal-fired boilers. The
first project, sponsored by EPRI, U. S. EPA, and Richmond Power and Light,
involves a 61 MM tangentially-fired boiler, Whitewater Valley No. 2 at
Richmond Power and Light. The objective of this project is to evaluate
different combinations of sorbents and coals (with varying sulfur contents),
to assess process performance variables and determine retrofit control
costs. Testing began in 1987 with sorbents including hydrated lime,
limestone, dolomite hydrated lime, and "promoted" lime sorbent (25).
Preliminary results of short-term tests indicate that sulfur dioxide
emissions can be reduced by 25-40% at calcium to sulfur ration of 2. A flue
gas humidification system was designed and installed to counteract
degradation of ESP performance resulting from use of sorbent injection.
The second demonstration is a 105 MW wall-fired unit, Edgewater No. 4,
owned and operated by Ohio Edison. This has been funded by the U. S. EPA,
DOE, and the Ohio coal Development Office. Testing that began in 1987
showed that the technology can exceed its sulfur dioxide control goal of 50%
at a calcium to sulfur ratio of 2 (26). The demonstration did encounter
problems with downstream particulate control because of the high resistivity
ash produced during sorbent injection. A full-scale flue gas humidification
system was added in 1988 to reduce ash resistivity and increase overall
sulfur dioxide removal.
The U. S. EPA and Combustion Engineering have begun the third
demonstration—a program to demonstrate furnace sorbent injection on a
tangential, coal-fired utility boiler at Virginia Power's Yorktown Unit No.
2. The overall objective is to achieve a significant reduction
(approximately 50%) in sulfur dioxide while minimizing any negative effects
on boiler performance during long-term operation of an integrated sorbent
injection system and a low-NO firing system (27). For moderate calcium to
sulfur ratios (2-3:1), S02 reductions of 40-53 percent were obtained using
calcium hydroxide and of 25-42 percent were obtained using limestone (28).
Two Canadian utilities (Saskatchewan Power and Ontario Hydro) also are
conducting full-scale furnace sorbent injection tests on lignite and
medium-sulfur bituminous coals, respectively. Using a limestone slurry, the
3-38
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Ontario Hydro demonstration system was able to remove up to 70% of the
sulfur dioxide at a calcium to sulfur ratio of 2.5 to 3.
Baltimore Gas and Electric, Atlantic Electric, the Electric Power
Research Institute, and Babcock and Wilcox have co-funded a program to
evaluate the potential of dry sorbent injection technology on
cyclone-equipped units (29). Upper-furnace sorbent injection for sulfur
dioxide capture has been examined using a 6 million Btu/hr cyclone-equipped
boiler simulator. The major operating problems which resulted was a
significant increase in opacity from 10 to 60 percent.
Applicability
FSI can be applied to both new and retrofit applications. However,.
under existing utility NSPS, application to new boilers will be limited to
units burning low-sulfur coals where 70 percent SO- removal is acceptable.
Retrofit of FSI technology into an existing boiler will require installation
of additional equipment as well as modifications to the boiler and particu-
late control device. Additional equipment includes sorbent preparation,
storage, handling, and injection systems; participate control modifications;
and a new or expanded ash handling system.
FSI can increase the particulate loading in the boiler and particulate
control device by two- to three-fold, depending on the Ca/S ratio and the
coal sulfur and ash content. The increase in particulate loading as a
function of Ca/S ratio and coal sulfur content for a typical 10 percent ash
coal .is shown in Figure 3.4-5.
Fouling and slagging of boiler tubes may become excessive due to
increased solids loading and, particularly for acidic coal ashes, reduced
ash fusion temperatures. These deposits reduce water wall heat transfer,
and as a result, the superheater and reheater steam temperature. Increased
dust loads may also increase erosion. Sootblower upgrading by increasing
the number of blowers, blowing pressure, and frequency should reduce the
severity of. heat transfer problems. Finned-tube economizers probably will
need to be replaced with larger bare-tube economizers with an in-line
arrangement. In some cases, derating of the boiler may be necessary even
after these modifications.
3-39
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a?
d>
'•6
w
•g
"5
CO
500
400 -
300-
200 -
100 -
CaJS
0.5
I
1.0
2.0
3.0
4.0
Coal Sulfur Content, %
FIGURE 3.4-5. Increase in solids loading as a function of coal sulfur
content and Ca/S ratio for a typical 10% ash coal (17).
FSI may also affect performance of the existing particulate control
device (generally an ESP) due to the increased solids loading, the high
alkalinity of the sorbent, and low S03 level in the flue gas (30).
Pilot-scale tests indicate that flue gas humidification can return ESP
performance to near pre-sorbent Injection levels for most ESP's (26). For
baghouses, additional bag area, more frequent cleaning, and additional
induced draft fan capacity may be needed to handle the increased particulate
loading and pressure drop. Also, the capacity of the ash handling system
may need to be increased to handle the extra solids. Because of the large
amount of unreacted lime present in the solids, conversion of wet ash
handling systems to dry handling will also be necessary.
On the positive side, the lower S02/S03 content of the flue gas and
increased alkali In the flue gas particulate may improve boiler efficiency
by allowing lower boiler exit temperatures without cold-end acid corrosion.
3-40
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Cost
Estimated total capital costs for retrofitting FSI on an existing power
plant are much lower than for conventional FGD technologies (31). Depending
on the design of the boiler and existing participate control system, total
capital costs range from $35-50/kW if the existing participate control
system can be used without significant upgrading. If the existing ESP must
be replaced by a new fabric filter (e.g., due to space constraints that
limit installation of additional ESP plate area), total capital costs can
approach $150/kW.
Because of FSI's relatively low captial cost, total busbar costs are
very sensitive to operating costs associated with sorbent purchase and
disposal. Major factors affecting these costs are the Ca/S ratio required
to achieve a given SO- reduction, coal sulfur content, and unit costs for
sorbent and waste disposal. Also, because of the relatively low capital
cost of FSI, the cost per ton of S02 removed is relatively uniform over a
wide range of percent removals and plant sizes.
The following table presents the range of values for capital cost,
annualized cost, and cost per ton of pollutant removed using IAPCS for FSI
with humidification and upgrading of the existing ESPs (SCA = 300). A
sulfur dioxide removal efficiency of 70% was assumed based on the use of
humidification to improve both ESP performance as well as S02 capture. The
lower costs are for a large unit burning low sulfur coal and the higher
costs are for a small unit burning high sulfur coal. Figures 3.4-6 through
3.4-8 show these costs as a function of boiler size and coal sulfur content.
Range
Capital Cost ($/kW) 23.6 to 81.4
Annualized Cost (mills/kWh) 3.6 to 21.8
Cost Per ton of S02 Removed ($/ton) 516.1 to 4,995
Appendix H contains tables of costs as a function of boiler size, capacity
factor, coal sulfur and SO- reduction.
3-41
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x
oc
Ul
a.
Capacity Factor = 50%
502 Removal = 70% with
Humidificotion
100
300
500
700
900
1.100
1,300
MW
Figure 3.4-6. Furnace Sorbent Injection - Capital Cost
Current 1988 Dollars
1% S
2% S.
o 3% 5
V7. 5
Copacity Factor = 50%
S02 Removal = 70% with
Humidification
100
300
500
700
900
1,100
1.300
MW
Figure 3.4-7. Furnace Sorbent Injection - Levelized Annual Cost
Current 1988 Dollars
3-42
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!-w
«3
0.0
Copocity Factor = 50%
S02 Removal = 70% with
Humldificotion
100
300
500
700
MW
900
1,300
Figure 3.4-8. Furnace Sorbent Injection - Cost Per Ton
of S02 Removed, Current 1988 Dollars
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3.5 LOW-TEMPERATURE SORBENT INJECTION
Description
Low-temperature sorbent injection (LTSI) involves injection of dry or
s lurried sorbents into the duct area between the air preheater and participate
control device. Several different process concepts have been proposed. These
include injecting dry sodium compounds, powdered lime hydrates in conjunction
with water or steam, and lime slurries. Injection locations range from
immediately following the air preheater to the inlet of the particulate control
device. Potential advantages of these processes include simplicity and low
capital cost of sorbent preparation and injection equipment, and adaptability
to retrofit applications.
Sodium-Based Sorbents--
In sodium-based systems, SO. is captured in a gas-solid reaction. Key to
this reaction is the thermal decomposition of the sorbent to form a porous,
high surface area matrix of reactive sodium carbonate (NaCO).. The optimum
injection temperature is 275-350°F, physically just downstream of the air
heater. Because of its ability to "popcorn" during heating, nahcolite (a
naturally occurring mineral containing 70-90 percent sodium bicarbonate--
NaHC03) has very good reactivity and sulfur reduction performance. Trona (a
naturally occurring form of sodium sesquicarbonate--Na2C03'NaHC03'2H20) is more
densely packed and provides less surface area for reaction with S02> but is
easier to handle. In general, 20-30 percent of SO. removal occurs in the duct
while 70-80 percent of S0« removal occurs in the particulate control system.
For this reason, a baghouse which allows considerable contact between the flue
gas and sorbent is preferred.
Sulfur dioxide removal . performance with nahcolite and trona during
full-scale tests by the Colorado Springs (Colorado) Department of Utilities is
shown in Figure 3.5-1. At SO. removals of up to 80 percent, nahcolite
utilization approached 100 percent. NOX reductions of up to 23 percent were
also obtained during these tests (32). Earlier tests with soda ash (NaOH) were
relatively ineffective.
3-44
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a
u
a
CVJ
a
u
u
a
u
a.
ALL LOAD CONDITIONS
COLORADO COAL
O SODIUM SeSOUlCARBONATE
• SODIUM BICARBONATE
0,0
Figure 3.5-1. S02 removal as a function of normalized
stoichiometric ratio (NSR). (32)
3-45
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Calcium-Based Sorbents--
Both dry and slurried calcium-based sorbents are being investigated. For
both sorbents, flue gas cooling to within 20-30°F of adiabatic saturation
is critical (approximately 125°F in a typical coal-fired boiler system). To
achieve these temperatures with dry sorbents, water and/or steam is injected
into the flue gas downstream of the air preheater. In pilot-scale tests at
these temperatures, SO. removal rates exceeding 75 percent have been achieved
at Ca/S ratios of 2 using conventionally hydrated lime (33,34). Depending on
the sorbent and flue gas temperature, 70-90 percent of SO* removal occurs
during the Initial few seconds while the sorbent Is suspended In the flue gas.
The remaining 10-30 percent occurs in the particulate control device. The
quantitative advantages of a baghouse versus an ESP for SO. control are not
well defined.
Significant similarities in process chemistry exist between conventional
spray drying and low-temperature sorbent injection. For slurried sorbents, the
major differences are the available residence time and the requirement for
rapid sorbent drying to prevent caking of wet sorbent on duct walls.
Conventional spray dryers have a reactor vessel residence time of 5-12 seconds,
whereas duct spray drying must accomplish particle drying and SO. reaction in
two seconds or less. As a result, proper choice of atomizer design, slurry
concentration, and slurry injection rate is essential.
The concentration of sorbent in the slurry is a major variable in removing
SO. from the flue gas. In conventional lime-based spray dryers, the sol Ids
content of the slurry ranges from 10-40 percent by weight (lime slurries cannot
be readily pumped at solids concentrations above 35 percent). To avoid caking
on duct walls, the moisture content of the dried droplet should be less than 20
percent. Although high solids concentration slurry will dry more rapidly and
have less impingement on duct surfaces, by drying faster the slurry quickly
loses its reactivity and has less time to remove SO. from the flue gas.
Testing will be required to determine the optimum slurry concentration and
injection rate for a given SO. concentration and duct geometry.
3-46
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Two basic atomizer designs are available: two-fluid and rotary. The
two-fluid atomizer can project a narrow cone of spray, thus reducing
impingement of slurry on the duct walls. However, an air compressor is needed
and will be a significant energy consumer. Also, early experience with
two-fluid systems in conventional spray dryers encountered problems with
atomizer erosion and plugging.
Rotary atomizers offer reliability, thorough mixing, and flexibility in
varying slurry feed rates without altering the droplet size distribution.
Rotary atomizers are the predominant atomization method used in conventional
spray drying. However, rotary atomizers offer less control over the spray
pattern and may be more prone to duct wall caking problems.
By using a dry sorbent, caking problems are reduced, but flue gas
humidification is required to achieve reasonable sorbent utilization. Dry
injection has two potential mechanical advantages over slurry injection.
First, the sorbent material is handled in a dry form rather than as a slurry,
thus avoiding the need for a slurry mixing and handling system. Second, by
separately feeding water and sorbent, operation of the system can be better
controlled than with a single feed system. This is especially significant with
high-sulfur coals where the quantities of calcium hydroxide required for
significant S02 removal (>50 percent) are large. In such a case, the quantity
of water required to maintain slurry rheology could result in excessive cooling
of the flue gas. This flexibility in sorbent/water injection also can
facilitate system adjustments required by fluctuations in boiler load.
Performance
Several in-duct sorbent injection processes are being developed. The
U. S. Department of Energy (DOE) Flue Gas Cleanup Program is sponsoring
development of low-cost emissions control technology that can be installed on
existing coal-fired power plants if acid rain legislation is enacted. The
emphasis of this program is on in-duct sulfur capture. In 1985, three projects
were undertaken to test duct injection technology on slipstreams from operating
coal-fired power plants.
One of the processes tested in the DOE program was the Hydrate Addition at
Low Temperature (HALT) process. In this process, calcium hydrate particles
3-47
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are Injected Into the flue gas stream before the gas Is cooled by
humidification with a fine spray of water. Sulfur dioxide removal efficiencies
ranged from 42 to 52% (35). Higher SOp reductions of 60-75% have been obtained
with corresponding 23 to 28% sorbent utilization, although wall wetting and
deposition problems were encountered (36). Another process being evaluated in
the DOE program is the Bechtel Confined Zone Dispersion (CZD) process which
captures sulfur dioxide by spraying a finely atomized slurry of hydrated lime
into the duct of a utility boiler between the air heater and the ESP. Two
reactive lime reagents were used for a pilot-scale test. Sulfur dioxide
removal efficiencies greater than 50% were achieved (Ca to S ratio of 1.5:1)
using either pressure hydrated dolomitic lime or calcltlc lime. A five-month,
large-scale test was carried out when the CZD process was retrofitted onto one
of two parallel flue ducts on a 140 MW unit. Wall wetting and deposition were
generally a serious problem.
The General Electric in-duct scrubbing (IDS) process also has been tested
at the pilot plant scale under DOE sponsorship. This type of scrubbing is
accomplished by injecting a slaked lime slurry with a rotary atomizer directly
into existing ductwork. Sulfur dioxide removal takes place in three areas: the
slurry injection zone, the evaporation zone, and the drift zone. Test results
indicated a lime utilization of 35% calculated for a calcium to sulfur ratio of
1.1 at an S02 removal efficiency of 39%.
Another type of Injection process being sponsored by'EPA is E-SOX> This
process appears to be especially appealing for retrofit applications since it
makes use of an existing ESP modified for injecting a lime slurry into the
first field of the ESP. The front end of an existing ESP is converted to a
spray chamber where the gaseous sulfur dioxide comes Into contact with lime
slurry droplets. The remaining ESP fields are modified with cold pipe
prechargers to bring the ESP particulate emissions back to current levels.
The Ohio Coal Development Office, U. S. EPA, Babcock & Wilcox, and Ohio
Edison are jointly sponsoring an evaluation of this process at a 5 MW pilot
plant; a 10-month field pilot test program at Ohio Edison's R. E. Burger
Station was scheduled to begin in early 1989 (37). Testing will focus on
demonstrating a minimum 50% sulfur dioxide removal rate, maintaining acceptable
ESP performance with a reduced collector area (to make room for the spray
chamber) and increased particulate loading, and showing that solid waste can be
disposed of safely and economically.
3-48
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Another variation of sorbent injection is a combination of furnace and
duct injection. Tampella's LIFAC process, combines in-furnace injection with a
subsequent vertical humidification reactor for activating the unreacted calcium
oxide with water, thus increasing limestone utilization. The process can
achieve an S02 removal efficiency of 80-85% with commercial grade pulverized
limestone at calcium to sulfur ratios of 2.0 or less. Two full-scale
applications of this technology exist in Finland, and pilot-plant tests reveal
that the process is effective with coals of sulfur contents from 0.5 to 4.3%.
The Advanced Silicate (ADVACATE) process is an in-duct dry injection
process that uses a highly reactive calcium silicate hydrate sorbent. The
large surface area and high moisture retention capability of the sorbent makes
it very reactive to sulfur dioxide and more effective than hydrated lime for
dry injection. The sorbent can be produced by mixing a calcium-containing
material (e.g., lime) and a silica-containing material (e.g., fly ash) in water
at elevated temperature and pressure. The calcium silicate hydrate remains a
free flowing dry material with a moisture content up to 35% (37). Pilot and
bench scale testing have achieved greater than 80% SO- removal at a calcium to
sulfur ratio of 2.5 (38). The unreacted lime and fly ash collected in the
first particulate control section is hydrated to produce the calcium silicates.
The ADVACATE process can be used without furnace injection of limestone by
adding lime directly to the hydrator. Preliminary pilot testing of the
ADVACATE process has shown a combined furnace/duct SO* removal efficiency
greater than 90% with a calcium to sulfur ratio near 1.0.
Applicability
Reflective of its development status, a number of unresolved issues remain
regarding retrofit of low-temperature sorbent injection:-flue gas temperature
and velocity limitations and their significance in boiler operations (e.g.,
load fluctuation); performance on boilers firing high-sulfur coals; method of
water injection and limitations on flue gas approach to saturation temperature;
sorbent distribution and residence time; sorbent utilization rates at various
stoichiometrlc ratios, injection temperatures, and residence times; choice of
and impact on the exiting particulate control system; potential for sorbent
fallout and caking in the duct and particulate control systems; and disposal of
solid wastes containing high levels of unreacted lime. Because of
3-49
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similarities, much of the data collected from earlier furnace sorbent injection
testing should be directly transferable.
Equipment for LTSI can potentially be installed without major alterations.
To achieve satisfactory contact between the sorbent and flue gas, the sorbent
injection point should be located at a point where the gas flow is reasonably
uniform. Assuming a typical flue gas velocity of 50 feet per second at full
load, 50 to 100 feet of straight duct downstream from the point of injection
will be required for drying slurried sorbents. An estimated 75 percent of
existing coal-fired power plants burning medium- and high-sulfur coals have
adequate duct lengths to meet this requirement (39).
Cost
Although both sodium and calcium compounds are can be used with LTSI,
uncertainties regarding the availability and cost of sodium compounds in the
eastern U. S. and the disposal of sodium wastes are expected to limit their
use. Therefore, cost estimates in this section apply only to lime-based
systems.
Preliminary estimates of direct capital cost for lime injection systems
(assuming the existing ESP can be used without significant modification) ranged
from $ll-16/kW for dry injection (33,40) to $27-32/kW for slurry injection
systems (based on comparison with spray dryer costs). Whether the capital cost
differences between dry and slurry injection systems (as well as between
furnace and low-temperature injection) are real or result from variations in
costing procedures used by different estimators Is uncertain. If significant
modifications to the particulate control system are required, capital costs can
increase several-fold.
Sorbent utilization rates achievable with LTSI are also uncertain due to
the lack of clear understanding of process chemistry and limited test data.
Estimates of Ca/S ratios required to achieve 50 percent SO- removal range from
0.75 to 2.5. As a result, significant uncertainty exists regarding operating
costs for LTSI. Based on I) the uncertainties in capital and operating costs
for LTSI and 2) general similarities in hardware for LTSI and FSI, it is likely
that the cost of the two basic approaches to SO* control are similar.
Therefore, the costs presented in Section 3.4 for FSI are likely to be
applicable to LTSI.
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3.6 REBURNING
Description
Reburning involves bypassing a portion of the fuel around the main
combustion zone and injecting it above the main burners to form a reducing zone
in which NO is converted to reduced nitrogen compounds. The overall process
is divided into three zones:
t Main Combustion Zone. Approximately 80-85 percent of the total fuel
to the boiler is combusted in this zone. The burners operate at
air-fuel ratios less than or equal to those typical of a normal
boiler (i.e., 10 percent excess air). No boiler modifications are
required in this zone.
t Reburning Zone. The remaining 15-20 percent of the fuel is injected
downstream of the main combustion zone to create a fuel-rich
reburning zone. Nitrogen oxides (NO ) produced in the main
combustion zone react with hydrocarbon radicals formed by partial
oxidation of the reburning fuel to form reduced nitrogen species such
as NH3 and HCN. A portion of these compounds are converted to N£.
The degree of NO reduction depends on overall air-fuel ratio
(typically 0.80 to 0.95), temperature, residence time, and
concentration of NOX in the reburning zone.
t Burnout Zone. Additional air is added in the cooler, upper furnace
to create overall fuel-lean conditions, and ensure complete oxidation
of the reburning fuel. Reduced nitrogen species formed in the
reburning zone are converted to either NO or H*'
Any fuel can be used as the reburning fuel. However, U. S. pilot studies
indicate that fuels with Tittle or no fuel bound nitrogen, such as natural gas,
can achieve greater NO reductions. These advantages are particularly
significant under operating conditions of lower NO levels exiting the main
combustion zone (less than 200 ppm) and short residence times. The volatility
3-51
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and low fuel nitrogen content of natural gas are the primary contributing
factors. SOg reduction also results from use of sulfur-free natural gas for
reburning. The potential exists to further reduce SO- emissions by combining
reburning with furnace sorbent injection.
Applicability
Because the main combustion zone of the furnace operates at normal
air-fuel ratios, gas reburning 1s applicable to a wide range of wall,
tangential, and cyclone-fired boilers. Boiler modifications for reburning
involve Installation of additional fuel injection burners and air ports above
the top or cyclone burner row. Since the main combustion zone operates
essentially under original design conditions, problems such as flame
impingement and poor carbon burnout are minimized.
For retrofit applications, adequate space (and residence time) between the
top burner row and the furnace exit must be available for the additional levels
of fuel and air injection. If adequate space 1s not available, a loss of NO
reduction performance and/or boiler derating at full load would likely be
incurred. Also, if natural gas is used for reburning, a natural gas supply
should already be available at or near the plant; otherwise, installation of a
natural gas supply pipeline may cause retrofit costs to be unreasonable.
Performance
The development of reburning technology has occurred primarily in the U.S.
and Japan, with the largest demonstrations occurring in Japan. In the U. S.,
the Gas Research Institute (GRI) has sponsored bench-scale (25 KW) and
pilot-scale (3 MM) furnace tests at Energy and Environmental Research
Corporation (EER) (34). Various fuels including natural gas, high-nitrogen
residual oil, and coal have been evaluated as reburning fuels. These tests
showed a NO reduction capability of 50-75 percent when using 15-20 percent
natural gas as the reburning fuel. Use of oil and coal as reburning fuels
resulted in less NO reduction and increased carbon content 1n flyash.
In the U. S., three commercial demonstration projects are currently
planned; two of the projects use natural gas as the reburn fuel and one will
3-52
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use coal as the reburn fuel. These projects are planned for start-up in the
early 1990's.
In Japan, boiler manufacturer Ishikawajima-Harima Heavy Industries (IHI)
has conducted pilot-scale tests as well as tests on two full-scale utility
boilers, a 600 MW and a 55 MM unit. These tests were conducted using fuel oil
and pulverized coal as the reburning fuels, respectively. The IHI pilot tests
showed reburning with pulverized coal could achieve 40-50 percent NO
reduction. However, in the two full-scale field tests, NO reductions were
less than 20-25 percent using oil or coal as the reburning fuel (41).
Another Japanese manufacturer, Mitsubishi Heavy Industries (MHI), has
conducted tests on a 125 MW oil-fired utility boiler. MHI has also over the
last 3 years conducted a full-scale demonstration on a 600 MW oil/coal dual
fuel-fired unit. Low-nitrogen oil was used as the reburning fuel in this
long-term demonstration, and achieved 50 percent NO reduction (42).
n
Cost
The range of cost estimates for natural gas reburning applied to
coal-fired utility boilers using the IAPCS cost model is presented below. The
lower costs are for a large unit with a low fuel price differential (FPD). The
higher costs are for a small unit with a high FPD. Fuel price differential is
the cost per million Btu difference between the current coal and the reburn
fuel. Figures 3.6-1 through 3.6-3 present capital cost, annualized cost, and
cost per ton of pollutant removed as functions of plant size and FPD. A NO
^
reduction of 60 percent was assumed.
Range
Capital Cost ($/kW) 6.8 to 34.2
Annualized Cost (mills/kWh) 1 to 16.7
Cost Per ton of NOX Removed ($/ton) 384.7 to 6,418.8
Appendix I contains tables of cost as a function of boiler size, capacity
factor, FPD and percent gas substitution.
3-53
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o
a
15 % Gas Substitution
Fuel Price Oifferentiol = 1 $/MM8tu
Fuel Price Differential = 2 $/MMBtu
25 % Gas Substitution
Fuel Price Differential = 1 $/MMBtu
Fuel Price Differential = 2 $/MMBtu
Capacity Factor = 50%
100
300
500
700
900
1.100
1,300
MW
Figure 3.6-1. Natural Gas Reburning - Capital Cost
Current 1988 Dollars
Jt
10 -
9 -
8 -
7 -
6 -
Legend
15 % Gas Substitution
Fuel Price Differential = 1. S/MMBtu
Fuel Price Differential = 2 $/MMBtu
25 % Gas Substitution
Fuel Price Differentiol = 1 $/MMBtu
Fuel Price Differentiol = 2 $/MM9tu
Capacity Factor = 50%
5 -
4 -
100
300
500
MW
700
900
1.100
1.300
Figure 3.6-2. Natural Gas Reburning - Levelized
Annual Cost, Current 1988 Dollars
3-54
-------
H.£ —
4 -
3.8 -
3.6 -1
3.4 -
3.2 -
^ 3 -
o e 2 8 -
a40 2 6 -*
£i 2-4 1
~ 2.2 -
2 -
1.8 -
1.6 -
1.4.-
1.2 -
1 -
— -— « » :
Legend
15 % Gas Substitution
• Fuel Price Differential = 1 $/MMBtu
o Fuel Price Differential = 2 $/MMBtu
25 % Gas Substitution
+ Fuel Price Differentiol = 1 $/MM8tu
A Fuel Price Differentiol = 2 $/MM8tu
Copocity Factor = 50%
ft •"•
" •-
h
i i i i i i I i i i i
100 300 500 700 900 1.100 1,300
MW
Figure 3.6-3. Natural Gas Reburning - Cost Per Ton
of NOx Removed, Current 1988 Dollars
3-55
-------
References
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Applications., and Economics. AP-3109, Electric Power Research Institute,
Palo Alto, California, 1983, 424 pp.
2. Rib, D. M. and D. R. Plumley. Experience at Cool Water with General
Electric Combined Cycle Equipment. Presented at the EPRI Coal
Gasification and Synthetic Fuels for Power Generation Conference,
San Francisco, California, 1985.
3. Matchak, T. A., A. D. Rao, V. Ramanathan and M. T. Sander. Cost and
Performance for Commercial Applications of Texaco-Based Gasification
Combined-Cycle Plants (Volume 1). AP-3486, Electric Power Research
Institute, Palo Alto, California, 1984, 112 pp.
4. McCarthy, C. B. and W. N. Clark. Integrated Gasification/Combined Cycle
(IGCC) Electrical Power Production -- A Rapidly Emerging Energy
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5. delaMora, J. A., et al. Evaluation of the British Gas Corporation/Lurgi
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AP-3980, Electric Power Research Institute, Palo Alto, California, 1985,
196 pp.
6. Hartman, J. J., T. A. Matchak, H. E. Sipe and M. Wu. Shell-Based
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8. Dawkins, R. P.,.et al. Cost and Performance of Kellogg Rust
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10. Pietruszkiewicz, J., et al. An Evaluation of Integrated-Gasification-
Combined-Cycle and Pulverized-Coal-Fired Steam Plants. AP-5950, Electric
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11. Yerushalmi, J. Circulating Fluidized Bed Boilers. In: Proceedings of
the 88th National Meeting of the AIChE (Volume 1), 1980, pp. 490-521.
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January 1986, p. 18.
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13. Update of Project Review Presentation by Bechtel Group, Inc. RP-1860-3,
Electric Power Research Institute, Palo Alto, California, 1983.
14. U. S. Department of Energy, Report to Congress on the Relationship Between
Projects Selected for the Clean Coal Technology Program and the
Recommendations of the Joint Report of the Special Envoys on Acid Rain.
DOE/FE-0072, 1986. 28 pp.
15. Coal and Syn Fuels Technology. DOE Announces Clean Coal Winners. July
28, 1986. p. 1.
16. Pope Engineers. Conversion of Coal-Fired Boilers to AFBC, A Statistical
Evaluation (Draft), pp. 14-18.
17. Miller, M. J. S0? and NO Retrofit Control Technologies Handbook.
CS-4277-SR, Electric Powe? Research Institute, Palo Alto, California,
1985, 366 pp.
18. Balling, L. and D. Hein. DeNO Catalytic Converters for Various Types of
Furnaces and Fuels, Development, Testing, Operation. In: Proceedings:
1989 Joint Symposium on Stationary Combustion NO Control, Vol. 2,
EPA-600/9-89-062b (NTIS PB89-220537), June 1989,xp. 7A-27.
19. Bauer, T. K. and R. G. Spendle. Selective Catalytic Reduction for
Coal-Fired Power Plants: Feasibility and Economics. CS-3603,
Electric Power Research Institute, Palo Alto, California, 1984, 164 pp.
20. McElroy, M. and G. Offen. In-Furnace SO- Control for Pulverized-Coal
Boilers. EPRI Journal, 10(7):53-55, 198S.
21. Makansi, J. Limestone Injection Achieves 50 Percent SO- Removal With
Minimal Side Effects. Power, 129(2):88-92, 1985. *
22. Chang, J. C. S. and C. B. Sedman. Scale-up Testing of the ADVACATE Damp
Solids Injection Process. In: Proceedings: First Combined FGD and Dry SO-
Control Symposium, Vol. 3, EPA-600/9-89-036c (NTIS PB 89-172175), March *
1989, p. 8-122.
23. Makansi, J. Understand System Effects When Evaluating Sorbent
Injection. Power, 129(6):35-39, 1985.
24. Kokkinos, A., D. C. Borio, R. W. Koucky, J. P. Clark, C. Y. Sun, and D. G.
Lachapelle. Boiler Design Criteria for Dry Sorbent SO Control
With Low - NO Burners. In: Proceedings: First Joint Symposium on Dry
SO- and Simultaneous SO,/NOV Control Technologies, Volume 2,
EPA-600/9-85-020b (NTIS^PB8§-232361), July 1985, p. 32-1.
25. England, G. C., B. A. Folsom, R. Payne, T. M. Sommer, M. W. McElroy, P. J.
Chappel, and I. A. Huffman. Prototype Evaluation of Sorbent Injection
with Humidification. In: Proceedings: First Combined FGD and Dry S0«
Control Symposium, Vol. 1, EPA-600/9-89-036a (NTIS PB89-172159), MarCh
1989, p. 4-15.
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26. Nolan, P. S., R. V. Hendriks, and N. Kresovlch. Operation of the
LIMB/Humidifier Demonstration at Edgewater. In: Proceedings: First
Combined FGD and Dry SO, Control Symposium, Vol. 1, EPA-600/9-89-036a
(NTIS PB89-172159), Marth 1989, p. 4-1.
27. Goglnenl, M. R., J. P. Clark, J. L. Marion, R. W. Koucky, D. K. Anderson,
A. F. Kwasnik, E. Gootzait, D. G. Lachapelle, and S. L. Rakes.
Development and Demonstration of Sorbent Injection for SO, Control on
Tangentially Coal-Fired Boilers. In: Proceedings: First Combined FGD and
Dry S07 Control Symposium, Vol. 1, EPA-600/9-89-036a (NTIS PB89-172159),
March 1989, p. 4-35.
28. Stuart-Sheppard, I. J. and C. J. Barnett. Retrofit Applications for
Control of SO- and NO . In: Proceedings of the 81st Annual Meeting of
APCA. Dallas, Texas, June 1988.
29. Farzan, H., et al. Pilot Evaluation of Reburning for Cyclone Boiler NO
Control. In: Proceedings: 1989 Joint Symposium on Stationary Combustion
NOV Control, Vol. 1, EPA-600/9-89-062a (NTIS PB89-220529), June 1989, p.
3-?. /
30. Dahlih, R. S., J. P. Gooch, and J. D. Kilgroe. Effects of Furnace Sorbent
Injection on Fly Ash Characteristics and Electrostatic Precipitator
Performance. In: Proceedings: First Joint Symposium on Dry SO, and
Simultaneous SO,/NOV Control Technologies, Volume 2, EPA-600/9-85-020b
(NTIS PB 85-232361)7 July 1985, p. 29-1.
31. Lachapelle, D. G., N. Kaplan, and J. Chappell. EPA's LIMB Cost
Model: Development and Comparative Case Studies. In: Proceedings:
First Joint Symposium on Dry SO, and Simultaneous S0,/N0 Control
Technologies, Volume 2, EPA-60079-85-020b (NTIS PB 85-232361), July 1985,
p. 35-1.
32. Ablin, D.W., et al. Full Scale Demonstration of Dry Sodium Injection
Flue Gas Desulfurization at City of Colorado Springs Ray D. Nixon Power
Plant. In: Proceedings: 1986 Joint Symposium on Dry SO, and
Simultaneous S0,/N0tf Control Technologies, Volume 2, EPA-600/9-86-029b
(NTIS PB87-120457),"October 1986, p. 48-1.
33. Babu, M., et al. Results of 1.0 Million Btu/hour Testing and Plans
for a 5 MW Pilot HALT Program for SO- Control. In: Proceedings of the
Third Annual Pittsburgh Coal Conference, Pittsburgh, Pennsylvania, 1986.
34. Abrams, J. Z., R. M. Sherwin, and G. H. Dyer. Partial FGD by Confined
Zone Dispersion of Pressure-Hydrated Lime. In: Proceedings of Coal
Technology '85, Pittsburgh, Pennsylvania, 1985. pp. 153-166.
35. Drummond, C. J., et al. Duct Injection Technologies for SO, Control.
In: Proceedings: First Combined FGD and Dry SO, Control Symposium, Vol. 3,
EPA-600/9-89-036C (NTIS PB89-172175), March 1989, p. 8-24.
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36. Tischer, R. and C. Drummond. Duct Injection Technologies for SO. Control.
In: Proceedings of the 15th Annual Biennial Low-Rank Fuels Symposium.
DOE/METC-90/6109. St. Paul, Minnesota, May 22-25, 1989.
37. Hovis, L. S., et al. E-SO Pilot Evaluation. In: Proceedings: First
Combined FGD and Dry SO, Control Symposium, Vol. 3, EPA-600/9-89-036c
.(NTIS PB89-172175), MarCh 1989, p. 8-177.
38. Jorgensen, C.f et al. Pilot Plant Evaluation of Post-Comt>ustion LIMB SO-
Capture. In: Proceedings: First Combined FGD and Dry S09 .Control
Symposium, Vol. 2, EPA-600/9-89-036b (NTIS PB89-172167;, (). 5-91.
39, Shilling, N. Z. In-Duct Application of Dry Flue Gas Desulfurization: A
Cost Effective Retrofit for Reduction of Sulfur Emissions. In:
Proceedings of the Second Annual Pittsburgh Coal Conference, Pittsburgh,
Pennsylvania, 1985. pp. 158-163.
40. Yoon, H., P. A. Ring, and F. P. Burke. Coolside S02 Abatement Technology:
1 MW Field Tests. In: Proceedings of Coal Technology 1985, Pittsburgh,
Pennsylvania, 1985. pp. 129-152.
41. Miyamae, S., et al. Evaluation of In-Furnace NO Reduction. In:
Proceedings: 1985 Symposium on Stationary Combustion NO Control, Volume
1, EPA-600/9-86-021a (NTIS PB 86-225042), July 1986, p. 24-1.
42. Murakami, N. Application of the MACT In-Furnace NO Removal Process
Coupled with a Low NO SGR Burner. In: Proceedings: 1985 Symposium on
Stationary Combustion NO Control, Volume 1, EPA-600/9-86-021a
(NTIS PB 86-225042), Jul? 1986, p. 32-1.
3-59
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SECTION 4
EMERGING TECHNOLOGIES
INTRODUCTION
This section reviews two areas of emerging technology which are currently
being researched and developed, but have not been demonstrated using full-scale
equipment. Depending on their economics and pollution control performance,
however, it is possible that these technologies could be in commercial use by
the mid- to late-1990's. These technologies are:
• Pre-Combustion Controls
- Advanced Coal Cleaning (Section 4.1)
• Post-Combustion Controls
- Advanced FGD (Section 4.2)
4-1
-------
4.1 ADVANCED COAL CLEANING
Advanced coal cleaning is divided into two primary areas: 1) advanced
physical cleaning and 2} chemical cleaning. Research is also underway in
biological cleaning processes, but is currently limited to bench-scale
experimentation. Evaluation of several advanced physical cleaning technolo-
gies 1s being conducted through a joint DOE/EPRI program at the EPRI Coal
Cleaning Test Facility (CCTF) 1n Homer City, Pennsylvania.
Description
Advanced Physical Coal Cleaning--
Advanced physical coal cleaning processes are being developed to increase
the removal of pyritlc sulfur and mineral matter from fine coal (<28 mesh);
organic sulfur is not affected. Although froth flotation and dense media
cyclones have been used commercially on fines between 28 and 100 mesh, these
techniques are currently ineffective at separating materials below 100 mesh
(referred to a "ultrafines"). Additionally, flotation is more effective at
removal of mineral matter than at removal of pyrite (1).
Research and development in advanced cleaning processes is concentrated in
four areas: 1) dense media cyclones, 2) froth flotation, 3) electrostatic
separation, and 4) selective coalescence (2). Most of these processes are
being developed to clean coal fines produced during crushing at a conventional
cleaning plant. Currently, these fines are generally either discarded or
recombined with the coarser cleaned product without being cleaned. Briefly,
these processes are:
t Dense media cyclones use heavy liquids such as fluorocarbon
refrigerants to separate coal from mineral matter based on
gravity differences.
• Froth flotation involves fine grinding and multistage flotation
to liberate sulfur and mineral matter from coal based on
differences in hydrophobic properties.
4-2
-------
• Electrostatic separation involves feeding dry, electrically charged
pulverized coal onto the surface of a rotating drum. Differences in
the dielectric properties of clean coal versus pyrite and ash cause
the impurities to separate.
t Small coal particles will selectively coalesce (i.e., agglomerate)
into larger particles in the presence of an appropriate medium. The
coalesced particles can then be separated from undesirable impurities
that do not coalesce. Research is underway on both the basic
physical mechanisms that are involved in this phenomenon and the use
of novel, non-aqueous media such as liquid carbon dioxide.
Chemical Coal Cleaning--
Chemical cleaning processes are being developed to remove both pyritic and
organic sulfur from coal. A variety of methods and chemical reagents have been
used in chemical cleaning of coal, including alkali displacement,
oxydesulfurization, and chlorinolysis. The major research and development
efforts currently underway involve caustic leaching, microwave treatment, and
partial hydrogenation.
• TRW's Gravimelt process is the primary caustic leach process under
development. Gravimelt exposes coal to a mixture of molten
potassium and sodium hydroxide at temperatures of 615-735°F.
Sulfur reacts with these alkalis to form sulfides. Residence time
is generally 2 to 3 hours. The cleaned coal floats to the
top of the molten caustic and liberated mineral matter. The clean
coal is then recovered by skimming and washed with water to remove
soluble contaminants and residual caustic. Additional ash removal
is achieved using an acid wash. Sulfur in the spent caustic is
converted to sulfuric acid, calcium sulfate, or elemental sulfur
in the regeneration section. The regenerated caustic is
reconcentrated and recycled (3).
• Microwave treatment is based on the tendency of water, caustic,
sulfur, and other ash components to absorb microwave radiation more
readily than the other coal constituents. When exposed to microwave
4-3
-------
radiation for less than a minute in the presence of potassium
hydroxide, sodium hydroxide, and water, a portion of the sulfur is
converted to alkali sulfides. The product coal is water washed to
remove-the sulfides and subsequently acid washed to dissolve the
converted ash.
• Sulfur removal can also be achieved by hydrogenation at mild
liquefaction conditions. Organic sulfur is converted to hydrogen
sulfide which can in turn be converted to elemental sulfur or
sulfuric acid. After separation of pyrite and insoluble mineral
matter, the final product is cooled to form a clean solid.
Biological Coal Cleaning--
Several microorganisms have been shown in laboratory and bench-scale tests
to promote removal of organic sulfur and pyrite occurring on the surface of
coal particles (4,5). Biological desulfurization is a low energy process, has
low operating costs, and does not reduce the heating value of the coal product.
Because the microorganisms attack the coal particle surface, crushing enhances
removal rates. Organic sulfur removals of 25-35 percent have been reported in
bench-scale tests with -60 mesh coals. Biological treatment could be
especially useful for coals containing very finely disseminated pyrite and
organic sulfur that is generally not removable by mechanical techniques.
Potential problems are relatively long bioprocessing times (days to weeks)
and the production of acidic, corrosive leaching solutions. However,
bioprocessing of stored coal may reduce the time-factor problem, if corrosion
problems can be overcome. Additionally, thermophilic microbial action
(occurring at >50°C) can accelerate processing with some bacteria.
Recent evidence shows that microorganisms can render pyrite hydrophilic in
minutes, thereby making it amenable to separation from coal by techniques such
as oil agglomeration. Development of such a treatment process would be
valuable in that pyrite separation could be accomplished in minutes compared to
several days required to remove comparable amounts of sulfur through bacterial
oxidation of pyrite to sulfate. Still to be determined are process parameters
and the biological mechanism.
4-4
-------
Because of the limited data available on the performance and economics of
biological treatment processes, no additional consideration of their potential
is presented in this report.
Applicability
Advanced coal cleaning processes can be applied to both new and retrofit
installations. However, since advanced physical coal cleaning cannot remove
organic sulfur, its primary benefit is on coals with a high percentage of
fine-grained pyritic sulfur and ash. Chemical cleaning can remove both pyritic
and organic sulfur as well as mineral matter and is therefore applicable to a
wider range of coals.
Performance
Pyrite in most coals generally accounts for 50-70 percent of the total
sulfur. Assuming advanced physical cleaning can remove 90 percent of the
pyritic sulfur, it would result in a 45-60 percent reduction in total sulfur.
Chemical cleaning can remove over 95 percent of both pyritic and organic
sulfur, and up to 99 percent of the mineral matter. Neither approach reduces
the nitrogen content of coal and, thus, they have no effect on NO emissions.
Costs
Advanced physical and chemical cleaning processes are both more expensive
per ton of product than conventional cleaning methods. They are therefore of
primary use in cleaning coal fines which cannot be handled by conventional
technologies or when additional sulfur removal is needed to meet more stringent
regulatory limits.
Depending on the specific process, capital cost estimates for advanced
physical coal cleaning equipment range from $22,000 to $82,000 per ton/hr of
input capacity; O&M cost estimates range from $2-11 per ton of coal processed
(6). The total capital cost of an integrated coal cleaning plant incorporating
4-5
-------
an advanced physical cleaning process can exceed.$120,000 per ton/hr of input
capacity. Typical O&M expenses for the total plant range from $4-8 per ton of
raw coal.
Table 4.1-1-reflects costs for a 450 ton per hour (input) integrated plant
in which the raw coal is crushed to -3/4 Inch (7). Conventional cleaning is
used on 28 mesh x 3/4-inch material and advanced processes are used on material
<28 mesh and on coarse middling products. Input to the advanced cleaning
system equals 40-50 percent of the total plant feed (I.e., 180-225 tons per
hour). Total capital cost ranged from $103,000-130,000 per ton/hr input
capacity. O&M ranged from $5.20-6.82 per Input ton. Coal sulfur content was
assumed to be 3.5 percent. Btu and weight recoveries ranged from 83.5-94
percent and 68-85 percent, respectively. Sulfur levels (in Ib SO^/mlllion Btu)
decreased by 34-78 percent. Because the least capital cost system may have
higher O&M or fuel loss costs (and vice versa), the lowest cost for the overall
system is greater than summation of individual cost components.
The capital and busbar costs from Table 4.1-1 are significantly higher
than for conventional physical cleaning (Section 2.3). The cost effectiveness
values are fairly similar, however, reflecting the higher level of sulfur
removal achievable. The cost for cleaning a 2 percent sulfur coal may be
similar, but due to the lower level of pyritic sulfur, the cost per ton of. SO-
removed may more than double.
The cost of chemical coal cleaning technologies is uncertain because of
the limited research results available (8). Current dollar cost estimates for.
the Gravimelt process presented In Table 4.1-2 range from $54-87 per ton of
cleaned coal (at 13,400 Btu/lb and 10 percent of the original sulfur). This
estimate is based on a 10,000 ton/day plant costing $270-380 million and
operating 330 days/year. As shown in Table 4.1-2, capital costs account for
20-25 percent of the annual 1 zed product price, O&M expenses 70-75 percent, and
fuel losses the remaining 5 percent or less.
These cost estimates do not include potential cost savings resulting from
improved boiler operations due to reduced solids loadings to the boiler and
emission control systems. These savings can significantly improve the
economics of coal cleaning, but are frequently specific to individual boilers
and coals.
4-6
-------
TABLE 4.1-1. ECONOMICS OF ADVANCED PHYSICAL COAL CLEANING
(3.5% Sulfur Coal)
$/Ton of Clean Coal
Capital
0 & M
Btu Loss
Total
Busbar Cost (mills/kWh)
$/ton of S0
Constant
2.70- 3.80
6.30- 9.00
2.10- 7.00
11.10-19.80
4.9 - 7.15
220-490
Current
5.20- 8.10
10.80-15.70
3.60-12.30
19.60-36.10
8.40-12.30
410-870
TABLE 4.1-2. ECONOMICS OF CHEMICAL COAL CLEANING
$/Ton of Clean Coal
Capital
0 & M
Btu Loss
Total
Busbar Cost (mills/kWh)
$/ton of SO-
Constant
6.00- 8.40
21.70-37.90
1.60- 2.20
29.30-48.50
11 -18
430-730
Current
12.50-17.60
37.90-66.60
2.70- 3.20
53.10-87.40
19 - 32
785-1300
4-7
-------
4.2
ADVANCED POST-COMBUSTION S02/NOX PROCESSES
Work 1s currently underway to develop improved methods for simultaneously
controlling S02 and NO emissions. Common elements in most of these processes
1s either reagent regeneration or production of a salable by-
product to reduce solid waste disposal volumes. This section focuses primarily
on two processes which are in pilot-scale and proof-of-concept testing:
electron-beam irradiation (E-Beam) and copper oxide. Other advanced processes
for simultaneous S02/NOX control which are less advanced include CONOSOX (based
on a potassium salt reagent) and Flakt Boliden (based on a sodium citrate
reagent).
Description
E-Beam--
Electron-beam processes Involve the irradiation of flue gas treated with a
reactant such as ammonia or lime to remove both SO- and NO. Factors
affecting S02 and NO removal by electron-beam irradiation include gas moisture
content, gas temperature, oxygen content, reagent ratio, and electron dosage.
In addition, efficient electron penetration of the gas stream requires a unique
discharge pattern and other special design considerations.
A simplified flow diagram of the E-beam/ammonia process is shown in
Figure 4.2-1. Incoming flue gas is cooled and humidified to about
10 percent moisture content in a water quench tower. The cooled flue gas is
injected with ammonia and then passed through an electron beam reactor. Oxygen
and water in the flue gas are ionized by the electrons to form the radicals
[HO], [0], and [H02] which react with S02 and NOX to form sulfuric acid (H2S04)
and nitric acid (HNO,). These acids are neutralized by the injected ammonia to
form solid ammonium sulfate ((NH.J-SO.) and ammonium sulfate nitrate
((NH.^SO '2NH.N03). Reaction time for formation of the sulfate and nitrate
salts is less than one second. Product solids are collected in the hopper
located below the reactor or in a downstream particulate collector.
4-8
-------
Ammonia
Quench Water
t t t t t
t t t t t
Flue Gas
E-Gun
1
E - Beam
Reactor
Drain
Participate
Collector
ID Booster
Fan
Product Solids
Figure 4.2-1. E-beam / ammonia process flow diagram (9).
-------
In another version of the E-beam process, the water quench tower and
ammonia injection system are replaced with a lime-based spray dryer. Reactions
are the same manner as above except that the products formed are calcium salts
(CaS04, Ca(N03)2, and CaSOj) instead of ammonium salts.
Copper Oxide--
Copper oxide (CuO) FGD is based on the reaction of copper oxide and SO- to
form copper sulfate (CuSO.). The CuSO,, and to a lesser extent CuO, catalyzes
the selective reduction of NO to N- in the presence of ammonia. Optimum
reaction temperature is 650-850°F. Spent CuSO. is sent to a second
vessel for regeneration by a reducing gas. The resulting concentrated S02
stream can be economically recovered as salable sulfur or sulfuric acid.
Two c.opper oxide processes have been developed. Shell Oil's version of
the process uses a set of specially designed, parallel-passage, fixed-bed
reactors containing copper oxide bonded to an alumina substrate. The other
approach, under development at DOE's Pittsburgh Energy Technology Center, uses
a fluidized bed of copper-impregnated alumina spheres.
A simplified diagram of the fluidized-bed process is shown in
Figure 4.2-2. In this configuration, the absorber is located between the plant
economizer and air heater; ammonia is injected upstream of the absorber. The
sulfated sorbent leaves the absorber through overflow pipes and weirs, and is
pneumatically transported with pre-heated air either to a solids heater or
directly to a moving-bed regenerator. The need for a solids heater depends on
the regeneration gas being used. In the regeneration vessel, the sulfated
sorbent contacts a reducing gas (CH4, H-, or a mixture of CO and H2) flowing
countercurrent to the sorbent. The regenerated sorbent leaves the bottom of
the reactor and is pneumatically conveyed in heated air to the absorber.
During transport, the copper formed during regeneration is oxidized to copper
oxide.
Other Advanced Concepts--
Durlng 1985, Argonne National Laboratory conducted laboratory and
bench-scale tests of two advanced SO^/NO processes. In laboratory tests
4-10
-------
To Air
Preheater
NH3
From
Economizer
Flue
I
Flue Gas
Gas
Fluidized-Bed
Absorber
Spent
Sorbent
Solids Heater
Grativating-Bed
Regenerator
CH4
To
Sulfur
Recovery
FIGURE 4.2-2. Simplified flow diagram for the Fluidized-Bed Copper Oxide
Process (10).
using double-alkali chemistry with proprietary additives, Argonne achieved
greater than 70 percent nitrogen oxides removal and 90 percent sulfur oxide
removal. Other laboratory work at Argonne has shown 50-70 percent nitrogen
oxide removal with greater than 90 percent S02 removal using lime spray drying
combined with additives; because this work applies to current FGD processes, it
could be scaled-up fairly quickly.
4-11
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ADD!icabllitv and Performance
E-Beam--
E-beam processes are in the early stages of development and have not been
tested at full-scale. Two major pilot projects are being conducted under the
U. S. Department of Energy Flue Gas Clean-up Program: one with Research-
Cottrell to evaluate the E-beam/lime spray dryer process, and the other with
Ebara International Corporation to evaluate E-beam/ammonia injection.
The E-beam/lime spray dryer pilot-scale system was installed on a
slipstream (4,000 ACFM) from a 150 MW coal-fired boiler at the TVA Shawnee
Steam Plant in Paducah, Kentucky. The pilot work began in early 1983 and was
completed in 1985. During the test program, five parameters were varied. The
inlet SO. concentration was observed to have the greatest impact on SO- and NO
£ t A
removal. At a high inlet SO- concentration (2,500 ppm) and an absorbed
electron dose of 1.5 Mrad (a Mrad equals 10 joules/g of flue gas), S02 removal
was greater than 82 percent and NO removal was greater than 90 percent. For a
low inlet SO- concentration (400 ppm), SO- removal was.greater than 95 percent
and NO removal was about 55 percent. The increased NO removal at high SO-
A AC
concentrations probably results from reaction of NO and NO- with H-SO. to form
HNOS04. (11)
The E-beam/ammonia Injection pilot-scale test facility is located at
Indiana Power and Light Company's E. W. Stout plant. Construction of the pilot
plant began in 1983. Testing completed in 1985 claims to have demonstrated 90
percent reduction in both SO- and NO . Detailed test results are not yet
available.
Copper Oxide--
The copper oxide process has been tested at both bench- and pilot-scale
levels. Bench-scale tests with a fluidized-bed absorber using flue gas from a
natural gas combustor doped to 3,000 ppmv have shown 90 to 95 percent SO-
removal and over 97 percent NO removal. Pilot-scale fluidized-bed tests using
flue gas from a 500 Ib/hour combustor burning a 3 percent sulfur bituminous
coal achieved 90 percent removal of both SO- and NO (12). In 1985, DOE
£ A
awarded a contract to scale this process up to the proof-of-concept level
4-12
-------
(i.e., 5 MW equivalent). This pilot-scale facility will be constructed and
operated at Commonwealth Edison's Kincaid Station (Kincaid, IL). The Shell
process has also been tested at the bench- and pilot-scale levels and has been
applied on a commercial 40 MW oil-fired boiler in Japan (13).
The integrated CuO process involves three major reactor systems (the
absorber, the regenerator, and SO- recovery) and a pneumatic solids handling
system between the absorber and regenerator. The relative complexity and space
requirements of this system plus the need for 650-850°F flue gas will limit the
potential for applying the CuO process to many existing boilers. One possible
alternative for overcoming space constraints is to install the absorber
downstream of the existing environmental control systems and to reheat the flue
gas to the desired temperature. However, operation at the lower end of the
desired temperature range (for economic reasons) will reduce S09/N0 removal
w A
rates. Because the sulfation reaction is exothermic, part of the reheat energy
may be recoverable from the cleaned flue gas.
Cost
Because of the relative complexity of these processes, retrofit onto
existing plants will be more difficult than other post-combustion FGD
technologies. The impact of this difficulty on retrofit costs has not been
quantified, however. Cost estimates for lime-based electron beam technology
are presented in Table 4.2-1. The costs shown for S02 control are the same as
for lime spray drying. The incremental costs for the system are included under
the costs for NOX control. The combined costs for the total system are shown
in the right-hand columns. Cost estimates for control of S02 and NOX using
copper oxide technology are presented in Table 4.2-2. In both cases, retrofit
factors for lime/limestone wet FGD were assumed.
4-13
-------
TABLE 4.2-1. COST ESTIMATES FOR ELECTRON BEAM
Technology
xso2
Removed 3.5X S
X NO
2X s Removed
S02 &
NO 3.5X S
X
NOx Combined
zx s
New 500-HW Baseload Power Plant:
Capital Cost, S/ku 70 - 90
MIlts/kUH, Current 70 - 90
S/Ton, Current S 70-90
Retrofit 500-HU Baseload Power Plant:
Capital Cost. S/kU 70 - 90
Mills/kWh, Current * 70-90
S/Ton. Current $ 70-90
Retrofit 250-NU Intermediate Load Power Plant:
Capital Cost, SAW 70 - 90 339 - 403
Mills/kwh. Current S 70-90 30.2 - 35.3
S/Ton. Current » 70-90 1.219 -1.609
174 - 207 153 - 182
15.7 - 18.6 12.0 • 14.0
677 - 804 880 - 1,081
297 - 352
25.2 - 29.2
1.753 - 2,365
90
90
90
226 - 270
17.1 - 20.4
730 - 883
198 - 235
13.2 - 15.4
961 - 1.202
90
90
90
90
90
90
93 - 109 284 - 315 248 - 275
6.3 - 6.7 22.4 - 25.2 18.3 - 20.1
2.632 - 3.732 930 - 1,135 1,298 • 1,629
122 - 143 369 - 408 323 - 357
7.0 - 7.6 24.7 - 27.8 20.3 • 22.4
2.944 - 4,182 1,017 -1,260 1,431 - 1,819
183 - 213 552 - 612 483 - 535
12.9 - 14.3 44.5 - 49.4 38.1 - 41.8
5,396 • 7,854 1,760 - 2.329 2,608 - 3,500
Retrofit factor of 1.0 and capacity factory of 65 percent.
Retrofit factor of 1.3 and capacity factor of 65 percent.
-------
TABLE 4.2-2. COST ESTIMATES FOR COPPER OXIDE
-e*
1-^
en
Technology Removed 3.5X S
X NO
2X S Removed NO
X
S02 &
3.5X S
NOx Combined
2% S
Neu 500-MU Baseloed Power Plant:
Capital Cost. S/kU 70 - 90
Mills/kUH, Current 70 - 90
S/Ton. Current $ 70-90
Retrofit 500-MU Baseload Power Plant:'
Capital Cost, J/kU 70 - 90
Hills/kUh. Current S 70-90
S/Ton, Current $ 70-90
187 - 203 155 - 167
17.7 - 20.5 13.0 - 14.7
760 - 883 949 • 1,143
235 - 255 192 - 207
18.8 - 21.9 14.0 - 15.8
813 - 951 1,028 - 1.236
Retrofit 250-MU Intermediate Load Power Plant:
Capital Cost, $/kW 70 - 90 360 - 391 290 - 315
Mills/kUh. Current $ 70-90 31.3 - 35.0 24.6 - 27.1
S/Ton. Current $ 70-90 1.269 -1.613 1.715 - 2,216
90
90
90
90
90
90
90
90
90
(10)
(12)
760 - 1,143
(10)
(12)
813 - 1.236
187 - 203 155 - 167
17.7 - 20.5 13.0 - 14.7
760 - 883 949 - 1.143
235 - 255 192 - 207
18.8 - 21.9 14.0 - 15.8
813 - 951 1.028 - 1,236
(10) 360 - 391 290 - 315
(12) 31.3 - 35.0 24.6 - 27.1
1.269 - 2.216 1.269 - 1.613 1.715 - 2.216
Retrofit factor of 1.0 and capacity factory of 65 percent.
Retrofit factor of 1.3 and capacity factor of 65 percent.
-------
References
1. Cavallaro, J. A., R. P. Killmeyer, A. W. Deurbrouck, and K. Rhee. An
Overview of PETC's Chemical, Physical, and Surface-Enhanced Beneficiation
Program. Pittsburgh Energy Technology Center, U. S. Department of Energy,
1986, 29 pp.
2. Engineering and Economics Research, Inc., et al. Supplemental Report to
Congress on Emerging Clean Coal Technologies. DOE/MC/22121-1. U. S.
Department of Energy, Washington, D.C., 1985.
3. Reference 2.
4. Olson, G. J., F. B. Brickman, and W. P. Iverson. Processing of Coal with
Microorganisms. AP-4472, Electric Power Research Institute, Palo Alto,
California, 1986, 48 pp.
5. Isbister, J. D., et al. Companion Processes for Removal of Sulfur and Ash
from Coal. In: Proceedings of the Second Annual Pittsburgh Coal
Conference, Pittsburgh, Pennsylvania, 1985, 809 pp.
6. Santhanam, C. J. and V. Vejins. Impact of Advanced Coal Benefication on
Utilization of Coal Slurry Fuels. In: Proceedings of the Seventh
International Symposium on Coal Slurry Fuels Preparation and Utilization,
New Orleans, Louisiana, 1985.
7. Boron, D. J., J. P. Hatoney, and M. C. Albrecht. Advanced Physical Coal
Cleaning: An Evaluation of Five Select Processes. In: Proceedings of
the Third Annual Pittsburgh Coal Conference, Pittsburgh, Pennsylvania,
1986, 974 pp.
8. Meyers, R. A., L. C. McClanathan, and W. D. Hart. Development of the TRW
Gravimelt Process. In: Proceedings of the Second Annual Pittsburgh Coal
Conference, Pittsburgh, Pennsylvania, 1985, 890 pp.
9. Aul, E., S. Margerum, and R. Berry. Industrial Boiler SO. Technology
Update Report. EPA-450/3-85-009 (NTIS PB85-197093), U.S.^Environmental
Protection Agency, Research Triangle Park, North Carolina, 1984.
10. Drummond, C. J., J. T. Yen, J. I. Joubert, and J. A. Ratafia-Brown. The
Design of a Dry Regenerative Fluidized-Bed Copper Oxide Process for the
Removal of Sulfur Dioxide and Nitrogen Oxides from Coal-Fired Boilers.
Presented at the 78th Annual Meeting and Exhibition of the Air Pollution
Control Association, 1985, 38 pp.
11. Gleason, R. J. and D. J. Helfritch. High-Efficiency NO and SO Removal
by Electron Beam. Chemical Engineering Progress, 81(10)?33-38, i985.
12. Reference 10.
13. Dickerman, J. C. and K. L. Johnson. Technology Assessment Report for
Industrial Boiler Applications: Flue Gas Desulfurization
EPA-600/7-79-1781 (NTIS PB 80-150873), U.S. Environmental Protection
Agency, Research Triangle Park, North Carolina, November 1979.
4-16
-------
APPENDIX A
Summary of Control Costs
Coal Switching and Blending
A-l
-------
COAL SWITCHING AI0 ILCNOINQ
CAPITAL COT (1/kU)
ESP FUEL Mice SULHA CAPACITY
SIZE DIFFERENTIAL (X) FACTOR
(«) («
LARGE ESP
(400 SCA)
LAIGB ESP
(400 SCA)
LARGE ESP
(400 SCA)
SMALL ESP
(200 SCA)
SMALL ESP
(200 SCA)
SMALL ESP
(200 SCA)
1.0
3 2.0 30.0 */W
3.0
4.0
1.0
5 2.0 SO.O */W
3.0
4.0
1.0
5 2.0 70.0 S/kW
3.0
4.0
1.0
S 2.0 30.0 */W
3.0
4.0
1.0
5 2.0 SO.O */kW
3.0
4.0
1.0
I 2.0 70.0 S/tt
3.0
4.0
BOILER GENERATING CAPACITY (NW)
100 300 500 700 1000
30.7
30.7
30.7
30.7
30.7
30.7
30.7
30.7
30.7
30.7
30.7
30.7
31.8
31.8
31.8
31. 8
31.8
31.8
31.8
31.8
31.6
31.8
31.8
31.4.
23.2
23. 2
23.2
23.2
23.2
23.2
23.2
23.2
23.2
23.2
23.2
23.2
24.3
24.3
24.3
24.3
24.3
24.3
24.3
24.3
24.3
24.3
24.3
24.3
21.3
21.3
21.3
21.3
21.3
21.3
21.3
21.3
21.3
21.3
21.3
21.3
22.4
22.4
ZZ.4
22.4
22.4
22.4
22.4
22.4
22.4
22.4
22.4
22.4
20.4
20.4
20.4
20.4
20.
20.
20.
20.
20.
20.
20.
20.
21.5
21. S
21.5
21.5
21.3
21.5
21.5
21. S
21.5
21.5
21.5
21.5
19.6
19.6
19.6
19.6
19.6
19.6
19.6
19.6
19.6
19.6
19.6
19.6
20.7
20.7
20.7
20.7
20.7
20.7
20.7
20.7
20.7
20.7
20.7
20.7
1300
19.1
19.1
19.1
19.1
19.1
19.1
19.1
19.1
19.1
19.1
19.1
19.1
20.2
20.2
20.2
20.2
20.2
20.2
20.2
20.2
43.1
4J.1
43.1
43.1
A-2
-------
COAL SWITCHING AND BLENDING
CAPITAL COST (S/kW)
ESP FUEL PRICE SULFUR CAPACITY
SIZE DIFFERENTIAL (%) FACTOR
(») . (X)
LARGE ESP
(400 SCA)
LARGE ESP
(400 SCA)
LARGE ESP
(400 SCA)
ouBBBonB BBBBSVI
SMALL ESP
(ZOO SCA)
SMALL ESP
(200 SCA)
SMALL ESP
(200 SCA)
1.0
15 2.0 50.0 S/ktt
3.0
4.0
1.0
15 2.0 50.0 S/kU
3.0
4.0
1.0
15 2.0 70.0 1/kW
3.0
4.0
raBBBBBBBBOBBBBiiBB*BBMBBnBABBBBBBBBBSttBWBBi
1.0
IS 2.0 30.0 t/W
3.0
4.0
1.0
IS 2.0 50.0 S/kW
3.0
4.0
1.0
15 2.0 70.0 S/kU
3.0
4.0
BOILER GENERATING CAPACITY (HW)
100 300 500 700 1000
41.1
41.1
41.1
41.1
41.1
41. 1
41.1
41.1
41.1
41.1
41.1
41.1
nmssvnsBi
42.2
42.2
42.2
42.2
42.2
42.2
42.2
42.2
42.2
42.2
42.2
42.2
33.6
33.6
33.6
33.6
33.3
33.5
33.3
33.5
33.3
33.3
33.5
33.3
•Bscnuu
34.7
34.7
34.7
34.7
34.7
34.7
34.7
34.7
34.7
34.7
34.7
34.7
31.6
31.6
31.6
31.6
31.6
31.6
31.6
31.6
31.6
31.6
31.6
31.6
32.8
32.8
32.8
32.8
32.7
32.7
32.7
32.7
32.7
32.7
32.7
32.7
30.7
30.7
30.7
30.7
30.7
30.7
30.7
30.7
30.7
30.7
30.7
30.7
31.8
31.8
31.8
31.8
31.8
31.8
31.8
31.8
31.8
31.8
31.8
31.8
29.9
29.9
29.9
29.9
29.9
29.9
29.9
29.9
29.9
29.9
29.9
29.9
tBBBBBQBVBB
31.1
31.1
31.1
31.1
31.1
31.1
31.1
31.1
31.1
31.1
31.1
31.1
1300
29.5
29.5
29.5
29.5
29.5
29.5
29.5
29.5
29.5
29.5
29.5
29.5
aaanas
30.6
30.6
30.6
30.6
30.6
30.6
30.6
30.6
30.6
30.6
30.6
30.6
A-3
-------
COAL SWITCHING AND BLENDING
ANNUAL COST (•Illl/Uh)
ESP FUEL PtlCB SUirm CAPACITY
SIZE DIFFERENTIAL (X) FACTM
(«> (W
LABCf f
-------
COAL SWITCHING AMD BLEND IKO
ANNUAL COST (•Ull/kWh)
ESP FUEL PRICE SULFUR CAPACITY BOILER GENERATING CAPACITY (MU)
SIZE DIFFERENTIAL (X) FACTOR
(<) (X) 100 300 500 700 1000
LARGE ESP
(400 SCA)
LARGE ESP
(400 SCA)
LARGE ESP
(400 SCA)
SMALL ESP
(200 SCA)
SMALL ESP
(200 SCA)
SMALL ESP
(200 SCA)
1.0
15 2.0
3.0
4.0
1.0
15 2.0
3.0
4.0
1.0
15 2.0
3.0
4.0
1.0
15 2.0
3.0
4.0
1.0
15 2.0
3.0
4.0
1.0
15 2.0
3.0
4.0
17.1
30.0 •UK/Ml 17.1
17.1
17.1
15.7
50.0 •UK/kUh 15.7
15.7
15.7
15.1
70.0 •UU/kUh 15.1
15.1
15.1
17.2
30.0 •
-------
COU. SWITCHING AMD BLENDING
$02 COST EFFECTIVENESS (t/TOH)
ESP FUEL PtICf SULFM CAPACITY
Size DIFFERENTIAL (I) FACTOR
(S) (X)
URGE ESP
(400 $CA)
UWGt ESP
(400 SOU
LARSE ESP.
(400 SCA)
SMALL ESP
(200 SCA)
SMALL ESP
(200 SCA)
SMALL ESP
(200 SCA)
1.0
5 2.0 30.0
3.0
4.0
1.0
S 2.0 SO.O
3.0
4.0
1.0
J 2.0 70. 0
3.0
4.0
1.0
S 2.0 30.0
3.0
4.0
1.0
3 2.0 50.0
3.0
4.0
1.0
I 2.0 70.0
3.0
4.0
BOILER GENERATING CAPACITY (NW)
100 300 500 TOO 1000
9650.8
S/TOH 916.8
481.2
326.3
8246.4
t/TOH 78S.3
412.2
279.1
7674
•/TON 729
182.7
259.4
970.8
I/TOI 929.4
487.9
330.8
8349.9
S/TOM 793.2
416.4
282.3
7736.4
I/TW 734.9
383.8
261. S
8147.4
774
406.3
273.4
7159.9
680.2
357
242
6736.9
640
333.9
227.7
8211.9
786.7
413
280
7244.3
668.2
361.2
244.9
6799.8
643.9
339.1
229.9
7797.3
740.7
388.8
263.6
6910
656.4
3U.6
233.6
6529.8
620.3
323.6
220.7
7931.9
753.3
395.5
268.1
6994.2
664.4
348.8
236.4
6592.5
626.3
328.7
222.9
7637.1
725. S
380.8
258.2
6797.9
645.8
339
229.8
6438.3
611.6
321.1
217.7
7771.7
738.3
387.5
262.7
6882.2
653.8
343.2
232.7
6501.1
617.6
324.2
219.8
7105.2
713
374.3
253.7
6706.4
637.1
334.4
226.7
6364
604.6
317.3
215.1
7640
725.8
381
258.3
6790.8
645.1
338.6
229.6
6426.8
610.3
320.3
217.3
1300
7428.7
703.7
370.4
251.1
66S3.3
632
331.8
224.9
6321.1
600.3
315.2
213.7
7543.3
718.5
377.2
255.7
6737.7
640
336
227.8
63B3.9
606.4
318.3
215. B
A-6
-------
COAL SWITCHING AND BLENDING
S02 COST EFFECTIVENESS (S/TON)
ESP FUE
SIZE OIF
LARGE ESP
(400 SCA)
LARGE ESP
(400 SCA)
LARGE ESP
(400 SCA)
SMALL ESP
(200 SCA)
SMALL ESP
(200 SCA)
SMALL ESP
(200 SCA)
L PRICE SULFUR
FERENT1AL (X)
<«).
1.
IS 2.
3.
4.
1.
15 2.
3.
4.
1.
15 2.
3.
4.
1.
15 2.
3.
4.
1.
15 2.
3.
4.
1.
15 2.
3.
4.
C
,0
,0
0
,0
0
,0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
APACITY
FACTOR
(X) 100
20220.6
30.0 «/TON 1920.9
1008.3
683.6
18510.4
50.0 S/TOV 1758.4
923
625.8
17777.6
70.0 S/TOM 1688.8
886.5
601
20354.1
30.0 S/TON 1933.5
1015
688.1
18593.9
50.0 S/TON 1766.3
927.2
628.6
17840
70.0 S/TON 1694.7
889.6
603.1
BOILER
300
18717.
1778.
933.
632.
17403.
1653.
667.
588.
16840.
1599.
839.
569.
18851.
1790.
940.
637.
17487.
1661.
87
591.
16903.
1605.
842.
571.
GE
4
1
4
8
3
2
a
3
3
7
8
3
9
8
1
3
6
2
2
2
2
7
9
4
NERATING
500
18367.4
1744.8
915.9
620.9
17153.5
1629.5
855.4
579.9
16633.3
1580.1
829.4
562.3
18501.9
1757.6
922.6
625.5
17237.7
1637.5
859.6
582.7
16696
1586
832.6
564.4
CAP AC M
700
18207.
1729.
907.
615.
17041.
1618.
849.
576.
16541.
1571.
824.
559.
18341.
1742.
914.
620.
17125.
1626.
as
578.
16604.
1577.
82
561.
rr
2
6
9
5
3
a
a
,1
a
4
9
2
8
4
6
1
6
8
*'
9
6
3
8
3
(MO
NX
18075.
1717.
901.
611.
16949.
1610.
845.
57
16467.
1564.
821.
556.
18210.
1729.
908.
615.
17034.
1618.
849.
575.
16530.
1570.
824.
558.
W
4
1
3
1
8
,1
2
"3
5
3
2
7
1
9
1
6
2
2
4
9
3
3
3
8
1300
17998.9
1709.8
897.5
608.5
16896. S
1605.1
842.6
571.2
16424.5
1560.2
819
555.2
18133.5
1722.6
904.2
613
16981.1
1613.1
846.8
574.1
16487.4
1566.2
822.2
557.4
A-7
-------
APPENDIX B
Low NOX Combustion Control Technologies
B-l
-------
NO* COMBUSTION CONTROL TECHNOLOGIES:
OVHRFIRI All, LOU Kb BLBNEI8
CAPITAL COST (S/kU>
FIRIMG
CONFIGURATION
TANGENTIAL
FIIING
(OvnriRE All)
TANCEITIAL
FIRING
(OVEIFIRC AIR)
TANGENTIAL
FIRING
(OVUFIIE AIR)
tMLL FIRING
(LOW NOB
BURNERS)
WALL FIIING
(Lay NOB
BURNERS)
UALL FIIING
(LOU NOx
HJBNIRS)
CAPACITY NOX
FACTOR REDUCTION
(I) «>
10
30 20 S/kW
50
10
50 20 »/kU
30
to
70 20 S/kv
30
20
30 40 S/kU
»
20
SO 40 S/kU
55
20
70 40 l/ky
51
BOILER GENERATING CAPACITY
100
.2
.2
.2
.
.
•
6.
6.
6.2
25.5
. 25.3
25.5
25.5
25.5
25.3
25.3
25.3
25.3
300
.2
.2
.2
.2
.2
.2
.2
.2
.2
13.2
13.2
13.2
13.2
13.2
13.2
13.2
13.2
13.2
500 700
2.4 .9
2.4 .9
2.4 .9
2.4 .9
2.4 .9
2.4 .9
2.4 .9
2.4 .9
2.4 .9
9.7 T.9
9.7 7.9
9.7 7.9
9.7 7.9
9.7 7.9
9.7 7.9
9.7 7.9
9.7 7.9
9.7 7.9
(NU)
1000
1.
1.
1. ~
1.
1.
1.
1.6
1.6
1.6
6.4
6.4
6.4
A.4
6.4
6.4
6.4
6.4
6.4
1300
1.
1.
1.
1.
1.
1.
1.3
1.3
1.3
5.5
3.3
5.5
S.S
5.5
5.5
5.5
5.5
5.5
B-2
-------
MOx COMBUSTION CONTROL TECHNOLOGIES:
OVERFIRE AIR, LOU NO* MINERS
ANNUAL COST (•itls/ktfl)
FIflINC CAPACITY NOx
CONFIGURATION FACTOR REDUCTION
(%) (X)
TANGENTIAL
FIRING 30
(OVERFIRE AIR)
TANGENTIAL
FIRING SO
(OVERFIRE AIR)
TANGENTIAL
FIRING 70
(OVERFIRE AIR)
VALL FIRING
(LOU NOx 30
BURNERS)
UALL FIRING
(LOU NOx SO
BURNERS)
UALL FIRING
(LOU NOx 70
BURNERS)
(X)
10
20
30
10
20
30
10
20
30
20
40
55
20
40
ss
20
40
SS
BOILER GENERATING CAPACITY (MU)
100
O.S
•UU/kUh O.S
O.S
0.3
•ilUAUh 0.3
0.3
0.2
nillt/kUh 0.2
0.2
2.1
Bllls/kUh 2.1
2.1
1-2
nllU/kWh 1.2
1.2
0.9
•ills/Ml 0.9
0.9
300
0.3
0.3
0.3
0.2
0.2
0.2
0.1
0.1
0.1
1.1
1.1
1.1
0.6
0.6
0.6
0.5
O.S
O.S
500
0.2
0.2
0.2
0.1
0.1
0.1
0.1
0.1
0.1
0.8
0.8
0.8
O.S
O.S
O.S
0.3
0.3
0.3
700
0.2
• 0.2
0.2
0.1
0.1
0-1
0.1
0.1
0.1
0.6
0.6
0.6
0.4
0.4
0.4
0.3
0.3
0.3
1000 1300
0.
0.
0.
0.
0.
0.
0.
0.
0.
0.1
0.1
0.1
0.1
0.1
0.1
0
0
0
O.S 0.4
O.S 0.4
0.5 0.4
0.3 0.
0.3 0.
0.3 0.
0.2 0.
0.2 0.
0.2 0.2
B-3
-------
NOX COMBUSTION COMTROl TfCNWLOOIES:
OVERFIRC All, LOU NO* UJUIRS
NO* COST EFFECTIVENESS (I/TON)
FIRING CAPACITY
CONFIGURATION FACTOR
(X)
TANGENTIAL
FIRING 30
(OVWItt AIR)
TANGENTIAL
MB INC SO
(OVERFIRE AIR)
TANGENTIAL
FIRING 70
{OVERFIRE AIR)
WALL FIRING
(LOU NO* 30
BURNERS)
MIL units
(LOU MX SO
BURNERS)
MALL FIRING
(LOU NO* 70
BUSKERS)
NOX
REDUCTION
(S)
10
20 I/TON
30
10
20 S/TON
30
10
20 I/TOM
30
20
40 I/TOM
IS
20
40 S/TON
55
20
40 S/TON
55
BOILER GENERATING CAPACITY (MU>
100
1621.2
810.6
540.4
972.7
466.4
324.2
694.8
347.4
231.6
2368.8
1194.4
868.6
1433.3
716.6
521.2
1023.8
511.9
372.3
300
838.8
419.4
279.6
503.3
251.6
167.6
359.5
179.7
119.8
1235.5
617.8
449. J
741.3
370.7
269.6
529.5
264. 8
192.6
500
617.3
308.7
205.6
370.4
185.2
123.5
264.6
132.3
68.2
909.4
454.7
330.7
545.7
272.6
196.4
389.8
194.9
141.7
700
504.7
252.4
168.2
302.8
151.4
100.9
216.3
108.2
72.1
743.1
371.6
270.2
445.9
222.9
162.1
318.3
159.2
115.6
1000
407.6
203.6
135.9
244.6
122.3
61 .5
174.7
87.4
56.2
600.1
300
218.2
360
180
130.9
257.2
126.6
93.5
1300
348
174
116
208.8
104.4
69.6
149.1
74.6
49.7
512.6
256.3
186.4
307.6
153.8
111.6
129.7
109.9
79.9
B-4
-------
APPENDIX C
Lime/Limestone FGD
C-l
-------
1/S-fCO
CAPITAL COST (»/kH)
SULFUI CAPACITY UTIOFIT
CONTENT . MCTOI FACTOt
(X) «> ' (X)
BQIUI GmfUTItt CAMCITY (NO
100 300 100 TOO 1000
1300
1.0 10
1.0 50
1.0 70
2.0 50
2.0 SO
2.0 70
3.0 30
3.C SO
3.0 70
4.0 30
4.0 SO
4.0 70
1.0
1.S
2.0
1.0
1.S
2.0
. 1.0
1.S
2.0
1.0
1.S
2.0
1.0
1.5
2.0
1.0
1.S
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
t.O
1.5
2.0
»/kU
t/ku
S/KU
»/kW
»/ku
s/ku
s/ku
»/Ui
t/kM
t/ku
i/fcy
•/ku
296.6
426.3
356
2M.a
426.5
556.2
297
426.7
556.4
316.4
455.2
594
316.6
453.4
594.2
316.7
435.5
SM.4
328.4
472.5
616.6
328.6
472.7
616.8
328.
472.
616.
334.
481.
628.1
335
481.7
628.3
339.2
481.9
428.5
190.5
271
351.5
190.6
271.1
351.6
190.7
271.2
351.7
196.8
279.8
362.7
196.9
279.9
362.8
197
280
362.9
204
289.9
375.8
204.1
290
376
204.2
290.1
376.1
209.8
298
386.2
209.9
298.1
386.3
210
298.2
386.4
158.2
223.6
289.1
158.3
223.7
289.2
158.4
223.8
289.2
164.8
232.8
300.9
164.9
232.9
301
164.9
233
301.1
175.2
247.8
320.3
175.1
247.9
320.6
175.4
248
320.7
180.9
253.6
330.3
181
255.7
330.4
181.1
255.8
330.5
140.3
198
255.4
140.6
198.1
255.5
140.7
198.1
255.6
146.3
206
265.7
146.4
206.1
265.8
146.4
206.1
265.8
155.1
218.4
281.7
155.1
218.5
281.8
155.2
218.5
281.9
160.1
223.4
290.6
160.2
225.4
290.7
160.3
225.5
290.7
136.7
192.2
247.7
136.7
192.3
247.8
136.8
192.3
247.8
144.5
203.3
262.1
144.6
203.4
262.2
144.7
203.5
262.3
150.1
211
271.9
150.2
211.1
272
150 J
211.2
272
154.1
217.5
280.1
154.9
217.6
280.2
155
217.6
280.1
125
175.3
225.5
125.1
175.3
225.6
125.1
175.4
225.6
132
185.2
238.3
132.1
185.2
238.4
132.2
185.3
238.4
137.3
192.7
247.9
137.6
192.8
248
137.7
192.9
248.1
143
200.2
257.5
143.1
200.3
257.5
143.2
200.4
257.6
C-2
-------
L/S-FGD
ANNUAL COST
SULFUR CAPACITT RETROFIT
CONTENT FACTOR - FACTOR
(X) (X)
BOILER GENERATING CAPACITY (NU)
100 300 500 700 1000
1300
1.0 30
1.0 50
1.0 70
2.0 30
2.0 50
2.0 70
3.0 30
3.0 50
3.0 70
4.0 30
4.0 50
4.0 70
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
45.6
•Ul«/Wh 61.4
77.1
30
•i lit/Ml 39.4
48.8
23.2
•Ult/kUh 29.9
36.7
49
•Ult/kUh 65.8
82.7
32.4
•Ult/kUh 42.3
52.6
25.2
•Ult/kUh 32.4
39.6
51.3
•Ult/kUh 68.8
86.3
34.2
•Ult/kUh 44.7
55.2
26.8
•Ult/kUh 34.3
41.8
53
•Ult/kUh 70.8
88.6
33.6
•Ult/kUh 46.3
57
28.1
•Ult/kUh 35.7
43.3
29
38.7
48.5
19.5
25.4
31.2
15.4
19.6
23.8
30.6
40.7
50.8
20.9
26.9
33
16.7
21
25.3
32.
42.
53.
22.
28.
34.
18
22.5
26.9
34
44.7
55.4
23.7
30.1
36.5
19.2
23.8
28.4
24.3
32.2
40.1
16.6
21.3
26.1
13.2
16.7
20.1
26
34.2
42.5
18
22.9
27.9
14.5
18.1
21.6
28.1
37
45.8
19.7
25
30.3
16
19.8
23.6
29.7
38.8
47.9
21
26.3
31.9
17.3
21.2
25.1
21.7
28.7
35.7
IS
19.2
23.4
12.1
15.1
18.1
23.3
30.6
37.8
^16.3
20.7
25
13.3
16.4
19.3
25.3
S3
40.7
17.9
22.5
27.1
14.7
18
21.3
26.8
34.8
42.7
19.2
24
28.7
16
19.3
22.7
20.9
27.7
34.4
14.3
18.5
22.6
11.7
14.6
17.5
22.8
30
37.1
16
20.3
24.6
13.1
16.1
19.2
24.4
31.8
39.2
17.3
21.8
26.2
14.3
17.3
20.6
25.9
33.5
41.1
18.6
23.2
27.7
15.5
18.7
22
19.4
25.
31.
13.
17.
20.
11
13.6
16.2
21.1
27.6
34
13
18.8
22.7
12.3
15.1
17.8
22.7
29.4
36.1
16.3
20.3
24.3
13.5
16.4
19.3
24.3
31.2
38.2
17.6
21.8
25.9
14.8
17.7
20.7
C-3
-------
l/S-MD
SOZ COST EFFECTIVIMSS (»/TOH)
SUfl* CAPACITY HTKVIT
CONTENT FACTOi FACTOB
U) (Z)
•QUO C£«UTKC CAPACITY (W)
100
300
300
700
1000
1100
1.0 30
1.0 50
1.0 70
2.0 30
2.0 SO
2.0 70
1.0 30
3.0 SO
3.0 70
4.0 30
4.0 50
4.0 70
1.0
1.3
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
l/TOi
S/TON
S/TON
S/TON
S/TON
S/TON
S/TON
S/TON
S/TON
S/TON
S/TON
S/TON
6282.2
8450.5
10618.6
4125.1
5426
6726.9
3189.5
4118.7
5047.9
3372.6
4333
5693. 3
2230.1
2926.4
3622.7
1733
2232.3
2729.7
2357.1
3160.2
3963.3
1570
2051.9
2533.7
1229.1
1373.3
1917.3
1825.9
2438.8
3051.8
1229.8
1593.6
1961.3
965.9
1228.6
1491.3
3988.9
5334.1
6679.3
2984.6
1491.7
4298.8
2122.3
2698.8
1275.3
2109.2
2802.7
1496.2
1438
1854.1
2270.2
1148.6
1445.9
1743.1
1487.7
1966.6
2443.3
1025.5
1312.8
1600.1
826.3
1031.5
1236.7
1171.3
1539.9
1908.6
815.8
1037
1258.2
662.6
820.6
978.6
3340
4434
5328.1
2280.2
2936.7
3593.1
1824.5
2293.4
2762.3
1787.2
2356.4
2925.6
1237.3
1578.8
1920.3
1000.8
1244.8
1488.7
1292
1696.9
2101.9
903.4
1146.4
1389.4
736.4
910
1083.3
1023.6
1333.9
1648.2
723.7
911.1
1098.5
594.8
728.7
862.5
2993.2
3953.7
4914.2
2064.2
2640.5
3216.9
1665.4
2077.1
2488.7
1607
2106
2605.1
1125.2
1424.6
1724
918.3
1132.2
1346.1
1162
1315
1867.9
822.7
1034.5
1246.2
677
828.3
979.6
923.8
1196.5
1469.5
661.8
829.4
989
549.3
666.2
783.1
2885.3
3813.3
4741.6
1993.9
2550.8
3107.7
1611.9
2009.7
2407.5
1571.4
2063
2554.6
1101.3
1396.3
1691.2
099.6
1110.3
1321.2
1120
1459.4
1798.8
795.5
999.2
1202.8
656.5
801.9
947.4
890.9
1152.7
1414.3
640.6
797.7
954.7
533.3
645.5
757.7
2668.1
3308.2
4348.4
1859.6
2363.7
2867.8
1513.6
1873.6
2233.7
1435.6
1899.7
2343.9
1029.8
1296.3
1562.8
847.5
1037.9
1228.2
1042.5
1350.1
1657.7
747.6
932.2
1116.8
621.4
753.3
885.1
835.7
1074.9
1314
606.5
749.9
893.4
508.3
610.8
713.3
C-4
-------
APPENDIX D
Integrated Gasification Combined Cycle
D-l
-------
ICCC-t/kU
SULFUR CAPACITY HEAT BOILER GENERATING CAPACITY (HU)
CONTENT FACTOR RATE
(X) (X) (Btu/kUh) 100 300 500 700 1000 1300
1.0 50 8000 */kW 2,373 1,849 1,652 1,536 1,423 1,345
10000 2,816 2,199 1,967 1,829 1,695 1,603
1.0 70 8000 SAW 2,372 1,848 1,652 1,535 1,422 1.344
10000 2,815 2,198 1,966 1,828 1,694 1.602
2.0 50 8000 $/kW 2,409 1,872 1,672 1,553 1,437 1.358
10000 2,857 2,226 1,989 1,848 1,712 1,618
2.0 70 8000 S/kU 2,408 1,871 1,671 1,552 1,436 1.358
10000 2,856 2,225 1,988 1,847 1,711 1,617
3.0 50 8000 S/kU 2,438 1,891 1,687 1,566 1,449 1,369
10000 2,890 2,248 2,007 1,864 1,725 1,631
3.0 70 8000 S/kW 2,437 1,890 1,686 1,566 1,448 1,368
10000 2,869 2,247 2,006 1,863 1,724 1,630
4.0 50 8000 S/kV 2,463 1,908 1,701 1,578 1.459 1,379
10000 2,919 2,267 2,022 1,878 1.737 1,642
4.0 70 8000 S/kV 2462 1,907 1,700 1,577 1.459 1.378
10000 2,918 2,266 2,021 1,877 1.736 1,641
0-2
-------
IGCC-mills/kUh
SULFUR
CONTENT
(%)
1.0
1.0
2.0
2.0
3.0
3.0
4.0
(.0
CAPACITY HEAT
FACTOR RATE
<%) (Btu/kUh)
50 8000 mills/kUh
10000
70 6000 mills/kUh
10000
50 BOOO miUs/Wh
10000
70 8000 mills/kWh
10000
50 8000 mills/kWh
10000
70 8000 millsAUh
10000
50 BOOO mills/kUh
10000
70 8000 mills/kWh
10000
100
171.3
201.4
130.7
154.4
172.9
203.4
131.8
155.6
174.1
204.8
132.7
156.4
175.2
206.0
133.4
157.3
BOILER
300
131.2
156.8
102.0
122.3
132.3
157.9
102.7
123.0
133.0
158.7
103.3
123.7
133.7
159.4
103.6
124.1
GENERATING
500
118.7
142.3
93.1
112.0
119.5
143.2
93.6
112.5
120.1
143.9
94.0
113.1
120.6
144.6
94.2
113.4
CAPACITY
700
111.7
134.2
88.0
106.2
112.4
134.9
88.4
106.8
112.9
135.6
88.7
107.1
113.2
136.0
88.9
107.5
(MW)
1000
105.2
126.7
83.3
100.8
105.7
127.2
83.7
101.2
106.1
127.8
83.1
101.5
106.4
128.1
84.0
101.7
1300
100.8
121.6
80.3
97.1
101.2
122.2
80.5
94.5
101.5
122.5
80.7
97.8
101.9
122.9
80.9
98.0
D-3
-------
APPENDIX E
Atmospheric Fluidized Bed
E-l
-------
AFBC-$/kU
SULFUR CAPACITY HEAT BOILER GENERATING CAPACITY (MU)
CONTENT FACTOR RATE
(%) (%) (Btu/kUh) 100 300 500 700 1000 1300
1.0 50 9000 $/kU 2,018 1,554 1,386 1,289 1,195 1,132
11000 2,346 1,8H 1,621 1,509 1,401 1,328
1.0 70 9000 t/kU 2,018 1,554 1,386 1,289 1,195 1,132
11000 2,346 1,814 1,621 1,509 1,401 1,328
2.0 50 9000 t/kU 2,031 1,562 1,393 1,295 1,200 1,137
11000 2,361 1,823 1,628 1,515 1,407 1,333
2.0 70 9000 $/kU 2,031 1,562 1,393 1,295 1,200 1,137
11000 2,361 1,823 1,628 1,515 1,407 1,333
3.0 SO 9000 S/kU 2,042 1,568 1,398 1,299 1.205 1,141
11000 2,372 1,830 1,634 1,521 1,412 1,338
3.0 70 9000 $/kU 2,042 1,568 1,398 1,299 1.205 1,141
11000 2,372 1,830 1,634 1,521 1,412 1,338
4.0 50 9000 */kW 2,050 1,574 1,403 1,304 1,209 1,145
11000 2,382 1,836 1,640 1,526 1,416 1,342
4.0 70 9000 S/kU 2,050 1,574 1,403 1,304 1.209 1,145
11000 2,382 1,836. 1,640 1,526 .1,416 1,342
E-2
-------
AFBC-mills/kUh
SULFUR CAPACITY HEAT
CONTENT FACTOR RATE
(%) (X) (Btu/kUh)
BOILER GENERATING CAPACITY (HU>
100 300 500 700
1000 1300
1.0
50 9000 mills/lcUh 148.9 118.8 108.7 102.9 97.3 93.8
11000 172.9 139.5 127.9 121.5 115.3 111.1
1.0
70
9000
11000
mills/kWh
116.4. 95.4
135.8 111.8
87.7
103.6
83.5
98.9
79.6
94.5
77.0
91.7
2.0
50
9000
11000
iniUs/kUh
151.7
176.2
121.3
142.5
110.0
130.7
105.2
124.3
99.8
118.1
96.1
113.9
2.0
70 9000 mills/tcWh 119.0 97.3 90.0 85.8 81.9 79.3
11000 138.8 114.6 106.4 101.7 97.3 94.3
3.0
50 9000 mills/kUh 154.2 123.6 113.2 107.5 102.0 98.4
11000 179.4 145.3 133.7 127.0 120.8 116.7
3.0
70 9000 mills/kWh 121.5 99.6 92.2 88.0 84.0 81.6
11000 141.8 117.4 109.2 104.5 100.0 97.0
4.0
50 9000 mills/kWh 156.8 126.0 115.5 109.7 104.3 100.6
11000 182.4 148.1 136.5 129.9 123.6 119.4
4.0
70 9000 mills/kUh 123.7 101.9 94.3 90.1 86.3 83.7
11000 144.6 120.2 111.8 107.1 102.7 99.8
E-3
-------
APPENDIX F
Lime Spray Drying
F-l
-------
LIME SPRAY DRYING WITH REUSE OF THE EXISTING ESPs
LSD+ESP (URGE ESP:400 SCA)
CAPITAL COST <$/kV)
SULFUR CAPACITY RETROFIT
CONTENT FACTOR FACTOR
(X) (X) (X)
1.0
1 30 1.5
2.0
1.0
1 50 1.5
2.0
1.0
1 70 1.5
2.0
1.0
2 50 1.5
2.0
1.0
2 50 1.5
2.0
1.0
2 70 1.5
2.0
1.0
3 30 1.5
2.0
1.0
3 50 1.5
2.0
1.0
3 70 1.5
2.0
1.0
4 30 1.5
2.0
1.0
4 50 1.5
2.0
1.0
4 70 1.5
2.0
BOILER GENERATING CAPACITY (MW)
100 300 500 700 1000
t/kU .
S/kU
S/kW
S/kU
S/kU
SAW
S/kU
S/kU
S/kU
S/kU
S/kU
t/ku
209.8
301.8
393.7
210.8
302.8
394.7
211.8
303.7
395.6
231.6
333.7
435.8
233.5
335.6
437.7
235.2
337.3
439.4
255.2
368.4
481.6
257.8
371.0
484.2
260.3
373.5
486.7
279.6
404.3
529.0
282.9
407.6
532.3
286.0
410.7
535.4
106.1
152.5
198.9
106.8
153.2
199.6
107.5
153.9
200.3
117.6
169.2
220.9
118.9
170.6
222.2
120.2
171.8
223.5
130.2
187.6
245.1
132.1
189.5
247.0
133.9
191.4
248.8
143.2
206.7
270.2
145.7
209.2
272.7
148.1
211.5
275.0
84.8
122.1
159.4
85.4
122.7
160.0
86.0
123.3
160.6
94.2
135.7
177.3
95.3
136.9
178.4
96.5
138.0
179.6
104.5
150.8
197.1
106.2
152.5
198.8
107.9
154.2
200.5
115.3
166.5
217.7
117.5
168.7
220.0
119.7
170.9
222.1
75.5
108.9
142.3
76.1
109.5
142.9
76.6
110.0
143.4
84.0
121.2
158.4
85.1
122.3
159.5
86.2
123.4
160.6
93.4
134.9
176.4
95.0
136.5
178.0
96.6
138.1
179.6
103.1
149.1
195.1
105.3
151.2
197.2
107.3
153.3
199.3
68.5
98.9
129.4
69.0
99.5
130.0
69.6
100.0
130.5
76.3
110.3
144.2
77.3
111.3
145.3
78.4
112.4
146.3
85.0
122.9
160.8
86.5
124.4
162.3
88.0
125.9
163.8
94.0
136.0
178.1
96.0
138.0
180.1
98.0
140.0
182.1
1300
64.7
93.6
122.5
65.3
94.2
.123.1
65.8
94.7
123.6
72.2
104.4
136.6
73.2
105.4
137.7
74.2
106.4
138.7
80.5
116.4
152.4
82.0
117.9
153.9
87.5
123.5
159.5
89.1
129.0
168.9
95.1
135.0
175.0
97.3
137.3
177.2
F-2
-------
LINE SPRAY DRYING WITH REUSE OF THE EXISTING ESPs
LSD+ESP (LARGE ESP:400 SCA)
ANNUAL COST (mills/kWh)
SULFUR CAPACITY RETROFIT
CONTENT FACTOR FACTOR
(%) (%)
1
1
1
2
2
2
3
3
3
4
4
4
30
50
70
30
50
70
30
50
•
70
30
50
70
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0.
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1-0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
miUs/kUh
(irills/kUh
mills/kUh
mills/kWh
mills/kUh
mills/kUh
mills/kUh
mills/kUh
mills/kWh
mills/kUh
mills/kUh
mills/kUh
BOILER GENERATING CAPACITY (MU)
100 300 500 700 1000
30
39.9
49.8
19.8
25.7
31.6
15.5
19.8
24
32.8
43.8
54.8
22
28.6
35.2
17.7
22.4
27.1
35.8
48
60.1
24.3
31.6
38.9
19.9
25.1
30.3
38.8
52.3
65.7
26.7
34.7
42.8
22.1
27.9
33.7
15.7
20.7
25.7
10..7
13.7
16.7
8.7
10.9
13
17.4
22.9
28.5
12.2
15.6
18.9
10.4
12.8
15.1
19.2
25.4
31.5
13.8
17.5
21.2
12
14.7
17.3
21
27.9
34.7
15.4
19.5
23.6
13.7
16.7
19.6
12.8
16.8
20.8
8.9
11.3
13.7
7.4
9.1
10.8
14.2
18.7
23.2
10.2
12.9
15.6
8.9
10.8
12.7
15.8
20.8
25.7
11.7
14.7
17.7
10.5
12.6
14.7
17.4
22.9
28.4
13.2
16.5
19.8
12
14.4
16.8
11.5
15.1
18.7
8.1
10.3
12.4
6.8
8.3
9.9
12.9
16.9
20.9
9.4
11.8
14.2
8.3
10
11.7
14.3
18.8
23.3
10.8
13.5
16.2
9.8
11.7
13.6
15.8
20.8
25.7
12.2
15.2
18.1
11.3
13.5
15.6
10.7
13.9
17.2
7.6
9.5
11.5
6.4
7.8
9.2
11.9
15.6
19.2
8.8
11
13.2
7.9
9.4
11
13.3
17.4
21.5
10.2
12.6
15.1
9.4
11.1
12.9
14.7
19.2
23.8
11.5
14.2
17
10.9
12.8
14.7
1300
10.3
13.4
16.5
7.4
9.2
11.1
6.3
7.7
9
11.5
15
18.4
8.6
10.7
12.8
7.7
9.2
10.7
12.8
16.7
20.6
9.9
12.3
14.6
9.4
11
12.7
14.2
18.5
22.8
11.5
14.1
16.7
10.9
12.7
14.6
F-3
-------
LIME SPRAT DRYING WITH REUSE OF THE EXISTING ESPS
LSD+ESP (LARGE ESP:400 SCA)
S02 COST EFFECTIVENESS (S/TOH)
SULFUR CAPACITY RETROFIT
CONTENT FACTOR FACTOR
(X) (X) (X)
1 30
1 50
1 70
2 . 30
,•{• ' ••".•1"
2 50
2 70
3 30
3 50
3 70
4 30
4 50
4 70
1.0
1.5
" 2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
• -.1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
1.0
1.5
2.0
S/TON
S/TON
S/TON
S/TON
S/TON
S/TON
S/TON
S/TON
S/TON
S/TON
S/TON
S/TON
BOILER GENERATING CAPACITY (HU)
100 300 500 700 1000
4868.4
6476.4
8084.6
3208.9
4173.7
5138.6
2526.2
3215.4
3904.6
2661.6
3554.6
4447.6
1784.5
2320.3
2856.1
1436.6
1819.3
2202
1936.6
2596.7
3256.7
1315.9
1711.9
2107.9
1077.7
1360.5
1643.4
1577.5
2122.9
2668.3
1083.5
1410.7
1737.9
899.5
1133.2
1366.9
2549.9
3361.7
4173.5
1738.3
2225.3
2712.4
1419.4
1767.3
2115.2
1411.3
1863
2314.6
993.1
1264.1
1535
842.2
1035.7
1229.3
1038.1
1373
1707.8
748.3
949.3
1150.2
652.3
795.8
939.3
853.6
1131.1
1408.7
627.2
793.7
960.2
558.1
677.1
796.1
2074.5
2727
3379.5
1439.4
1830.9
2222.4
1196.1
1475.8
1755.4
1155
1518.4
1881.8
832
1050
1268.1
721.9
877.6
1033.4
854.1
1123.9
1393.7
632.8
794.7
956.6
566.1
681.7
797.3
705.4
929.4
1153.4
534.2
668.6
803
488.8
584.8
680.8
1872.7
2457
3041.3
1314.6
1665.2
2015.8
1104.3
1354.7
1605.1
1046.2
1371.7
1697.3
764.6
960
1155.3
672.3
811.8
951.3
775.8
1017.8
1259.8
584.3
729.5
874.7
530.4
634.1
737.8
642.4
843.5
1044.5
495.2
615.8
736.4
460.1
546.3
. 632.4
1730.5
2263.5
2796.6
1230.5
1550.3
1870.2
1045
1273.5
1501.9
969.1
1266.3
1563.5
718.9
897.2
1075.5
639.9
767.3
894.6
720.3
941.4
1162.4
551.2
683.9
816.5
506.9
601.7
696.4
597.5
781.4
965.2
468.4
578.7
689
441.1
519.9
598.6
1300
1667.8
2173.3
2678.8
1198.8
1502.1
1805.4
1026.7
1243.3
1460
934.7
1216.7
1498.6
701.2
870.3
1039.5
629.4
750.2
871
695.2
905
1114.8
538.1
664
789.9
507.9
597.9
687.8
577
751.6
926.2
466.9
571.6
676.4
441.5
516.4
591.2
F-4
-------
LINE SPRAY DRYING WITH A NEW BAGHOUSE
LSD+FF
CAPITAL COST ($/kU)
SULFUR CAPACITY RETROFIT
CONTENT FACTOR FACTOR
-------
LIME SPRAT DRYING WITH A NEW BAGHOUSE
LSD+FF
ANNUAL COST (mills/kUh)
SULFUR CAPACITY RETROFIT
CONTENT FACTOR FACTOR
(*) (X) (X)
1.0
1 30 "l.S
2.0
1.0
1 50 1.5
2.0
1.0
1 70 1.5
2.0
vo
2 30 1.5
2.0
1.0
2 50 1.5
2.0
1.0
2 70 1.5
2.0
1.0
3 30 1.5
2.0
1.0
3 50 1.5
2.0
1.0
3 70 1.5
2.0
1-0
4 30 1.5
2.0
1.0
4 50 1.5
2.0
1.0
4 70 1.5
2.0
BOILER GENERATING CAPACITY (MU)
100 300 500 700 1000
mills/kUh
mills/kUh
mills/kWh
mills/kUh
mills/kUh
mills/kUh
mllls/kuh
mills/kWh
mills/kUh
mills/kWh
mills/kUh
mills/kUh
40.2
50.1
59.9
26.1
32.0
37.9
20.2
24.5
28.7
43.0
53.9
64.9
26.3
34.8
41.4
22.3
27.0
31.7
45.9
58.1
70.2
30.6
37.8
45.1
24.5
29.7
34.9
49.0
62.4
75.7
32.9
40.9
49.0
26.7
32.5
36.2
24.0
19.0
34.0
15.9
18.9
21.9
12.6
14.7
16.9
25.7
31.2
36.8
17.4
20.7
24.0
14.2
16.6
19.0
27.4
33.6
39.8
19.0
22.7
26.4
15.9
18.5
21.1
29.3
36.1
42.9
20.6
24.6
28.7
17.5
20.4
23.4
20.7
24.7
28.7
13.8
16.2
18.6
11.1
12.8
14.5
22.1
26.6
31.0
15.2
17.8
20.5
12.6
14.5
16.4
23.7
28.6
33.6
16.6
19.6
22.5
14.1
16.2
18.3
25.2
30.7
36.2
18.0
21.3
24.6
15.6
18.0
20.4
19.7
23.3
26.8
13.2
15.3
17.5
10.6
12.1
13.7
21.0
25.0
29.0
14.5
16.9
19.3
12.0
13.7
15.5
22.5
26.9
31.4
15.8
18.5
21.2
13.5
15.4
17.4
23.9
28.9
33.8
17.2
20.2
23.1
15.0
17.2
19.3
18.6
21.9
25.1
12.5
14.5
16.5
10.1
11.5
12.9
19.9
23.5
27.2
13.8
16.0
18.2
11.8
13.4
14.9
21.2
25.3
29.4
15.5
17.9
20.3
13.3
15.0
16.8
22.6
27.1
31.6
16.8
19.5
22.2
14.8
16.7
18.6
1300
18.5
21.6
24.7
12.6
14.4
16.3
10.2
11.5
12.8
19.8
23.2
26.7
13.8
15.9
17.9
11.6
13.1
14.6
21.1
25. 0
28.8
15.1
17.4
19.7
13.0
14.7
16.3
22.5
26.8
31.0
16.4
19.0
21.5
14.5
16.3
18.2
F-6
-------
LIME SPRAY DRYING WITH A NEW BAGKOUSE
LSD+FF
S02 COST EFFECTIVENESS (S/TON)
SULFUR CAPACITY RETROFIT
CONTENT FACTOR FACTOR
(X) (X) (X)
1
1
1
2
2
2
3
3
3
4
4
4
1.0
30 1.5
2.0
1.0
50 1.5
2.0
1.0
70 1.5
2.0
1.0
30 1.5
2.0
1.0
50 1.5
2.0
1.0
70 1.5
2.0
1.0
30 1.5
2.0
1.0
50 1.5
2.0
1.0
70 1.5
2.0
1.0
30 1.5
2.0
1.0
50 1.5
2.0
1.0
70 1.5
2.0
S/TON
S/TON
S/TON
S/TON
S/TON
S/TON
S/TON
S/TON
S/TON
S/TON
S/TON
S/TON
BOILER GENERATING CAPACITY
100 300 500 700
5790.9
7211.6
8632.4
3757.0
4609.5
5462.0
2913.5
3522.3
4131.3
. 3094.5
3883.5
4672.4
2035.5
2508.9
2982.3
1607.6
1945.9
2284.1
2205.3
2788.5
3371.7
1467.4
1817.3
2167.2
1176.9
1426.9
1676.8
1763.8
2245.7
2727.5
1185.3
1474.5
1763.6
962.8
1169.3
1375.8
3460.8
4177.9
4895.0
2289.8
2720.1
3150.3
1816.3
2123.7
2431.0
1849.4
2248.4
2647.4
1252.7
1492.1
1731.4
1023.2
1194.2
1365.2
1318.1
1613.9
1909.7
910.4
1087.9
1265.4
761.2
888.0
1014.7
1054.2
1299.5
1544.7
740.3
887.4
1034.6
630.9
736.0
841.1
2981.6
3558.0
4134.3
1990.1
2335.9
2681.7
1593.5
1840.5
2087.5
1593.7
1914.6
2235.6
1092.7
1285.3
1477.9
904.3
1041.8
1179.4
1136.0
1374.3
1612.7
796.5
939.6
1082.6
676.6
778.8
880.9
908.8
1106.7
1304.6
649.4
768.2
886.9
563.5
648.3
733.1
2834.1
3350.2
3866.2
1898.2
2207.9
2517.5
1525.4
1746.6
1967.8
1513.0
1800.5
2088.1
1042.5
1215.0
1387.5
867.1
990.3
1113.5
1078.1
1291.8
1505.6
760.5
888.7
1017.0
649.9
741.5
833.1
861.5
1039.2
1216.8
620.0
726.6
833.2
541.7
617.8
693.9
(MU)
1000
2681.0
3151.8
3622.6
1807.3
2089.8
2372.3
1461.2
1663.0
1864.8
1431.3
1693.7
1956.2
993.8
1151.3
1308.8
850.4
962.9
1075.4
1019.1
1214.4
1409.6
742.0
859.2
976.3
637.4
721.0
804.7
814.5
976.9
1139.3
604.9
702.3
799.7
531.4
601.0
670.6
1300
2672.0
3118.5
3564.9
1807.9
2075.8
2343.7
1465.9
1657.3
1848.6
1425.0
1674.0
1922.9
993.4
1142.8
1292.2
834.6
941.3
1048.0
1013.6
1198.9
1384.2
724.4
835.6
946.8
626.0
705.4
784.8
809.3
963.5
1117.7
590.7
683.2
775.8
522.2
588.3
654.4
F-7
-------
APPENDIX G
Selective Catalytic Reduction
G-l
-------
POST-CWBUrriW NOB aMTROlS:
SELECTIVE CATALYTIC tEDUCTKM
CAPITAL COST (S/kU)
CATALYST Ml CAPACITY 1ET80MT
lift REDUCTION f ACTOR r ACTOR
(TEARS) (X) (X) (X)
BOILER GENERATINO CAPACITY (MU)
100 500 500 700 1000 1300
1.00
1 70 30 f.SO
2.00
1.00
1 80 30 1.50
2.00
1.00
1 70 50 1.SO
2.00
1.00
1 80 50 1.30
2.00
1.00
1 70 70 1.50
2.00
1.00
1 80 70 1.50
2.00
1.00
3 70 30 1.50
2.00
1.00
3 80 30 1.50
2.00
1.00
3 70 SO 1.50
2.00
1.00
3 80 SO 1.50
2.00
1.00
3 70 70 1.50
2.00
1.00
3 SO 70 1.50
2.00
1/ku
SAW
t/ut
s/KU
S/kW
s/kv
t/kw
t/ku
»/fcW
>/kw
*/ku
»/ku
133
176.7
220.5
143
191.8
240.}
133
176.8
220.5
143.1
191.8
240.6
133
176.8
220.6
143.1
191.9
240.6
133
176.7
220.5
143
191.8
240.5
133
176.8
220.5
143.1
191.8
240.6
133
176.8
220.6
143.1
191.9
240.6
103.2
133
162.8
111
144.7
178.4
103.2
133
162.8
111.1
144.8
178.5
103.2
133
162.8
111.1
144.8
178.5
103.2
133
162.8
111
144.7
178.4
103.2
133
16T8
111.1
144.8
178.5
103.2
133
162.8
111.1
144.8
178.5
91.3
118.5
143.5
98.9
129.6
160.3
91.5
118.5
145.5
99
129.6
160.3
91.6
118.6
145.6
99
129.7
160.4
91.5
118.5
145.5
98.9
129.6
160.3
91. S
118.3
145.5
98.9
129.6
160.3
91. S
118.6
141.6
99
129.7
160.4
88.6
114.4
140.3
95.9
125.3
154.7
88.7
114.5
140.3
95.9
125.3
154.7
88.7
114.5
140.3
96
125.4
154.8
88.6
114.4
140.3
95.9
12S.3
154.7
88.7
114.1
140.3
95.9
125.3
154.7
88.7
114.5
140.3
96
129.4
154.8
86.6
111.5
136.4
93.7
122.1
150.6
86.6
111.6
136.5
93.8
122.2
150.6
86.7
111.6
136.5
93.8
122.2
150.7
86.6
111.5
136.4
93.7
122.1
150.6
86.6
111.6
136.5
93.7
122.2
150.6
86.7
111.6
136.3
93.8
122.2
150.7
85. 4
109.8
134.2
92.4
120.3
148.2
85.4
109.9
134.3
92.5
120.4
148.3
85.3
109.9
134.3
92.5
120.4
148.3
85.4
109.8
134.2
92.4
120.3
148.2
85.4
109.8
134.3
92.5
120.4
148.3
85.5
109.9
134.3
92.S
120.4
148.3
G-2
-------
POST'OMUSTION MR CWTKXS:
SELECTtVf CATALYTIC REDUCTION
ANNUM. COST (•fllt/lrwri)
CATALYST NO* CAPACITY RETROFIT
LIFE REDUCTION FACTOR FACTO*
(TEAR!) (X) (X) (X)
1.00
1 70 30 1.50
2.00
1.00
1 80 30 1.50
2.00
1.00
1 70 50 1.50
2.00
1.00
1 60 50 1.50
2.00
1.00
1 70 70-1.50
2.00
1.00
1 80 70 1.50
2.00
1.00
3 70 30 1.50
2.00
1.00
3 80 30 1.50
2.00
1.00
3 70 50 1.50
2.00
1.00
3 80 50 1.50
2.00
+
1.00
3 70 70 1.50
2.00
1.00
3 80 70 1.50
2.00
•OILER GENERATING CAPACITY (MW)
100 300 500 700 1000
•fttlAUl
•UU/tMh
•ttU/kUi
•Uti/ktfi
•UU/Mft
•UU/Uft
•ills/Mi
•lUl/blil
•UU/Ml
•llU/Uh
•iUs/kuh
•Ult/Mft
29.4
33.4
37.5
30.4
34.9
39.4
17.9
20.3
22.8
18.3
21.2
23.9
13
14.7
16.5
13.4
15.4
17.3
18.4
22.3
26.6
19.4
23.9
28.5
11.3
13.7
U.2
11.9
14.6
17.4
8.3
10
11.8
8.7
10.7
12.6
26.1
28.8
31.6
26.9
30
33.1
15.9
17.6
19.2
16.4
18.3
20.2
11.6
12.7
13.9
11.9
13.3
14.6
15.1
17.9
20.7
13.9
19.1
22.2
9.3
,11
ri-7
9.8
11.7
13.6
6.9
8
9.2
7.2
8.6
9.9
25
27.5
30
25.8
28.6
31.3
15.3
16.8
18.3
15.7
17.4
19.2
11.1
12.2
13.3
11.5
12.7
13.9
14.1
16.6
19.1
14.8
17.7
20.5
8.7
10.2
11.7
9.2
10.9
12.6
6.4
7.5
8.6
6.8
8
9.2
24.7
27.1
29.5
25.4
28.2
30.9
15.1
16.5
17.9
15.5
17.2
18.8
11
12
13
11.3
12.5
13.6
13.7
16.1
18.5
14.5
17.2
19.9
8.5
9.9
11.4
9
10.6
12.2
6.)
7.3
8.3
6.6
7.8
9
24.5
26.8
29.1
25.2
27.9
30.5
14.9
16.3
17.7
15.4
17
18.6
10.9
11.9
12.9
11.2
12.3
13.5
13.6
13.9
18.2
14.3
16.9
19.6
8.4
9.8
11.2
8.8
10.4
12
6.
7.
8.
6.
7.
8.8
1300
24.3
26.6
28.9
25.1
27.6
30.2
14.9
16.2
17.6
13.3
16.9
18.4
10.8
11.8
12.8
11.1
12.3
13.4
13.4
13.7
17.9
14.1
16.7
19.3
8.3
9.7
11
8.7
10.3
11.9
6.1
7.1
8.1
6.5
7.6
8.7
G-3
-------
MR CMTMX.S:
SCLECTIVI CATALYTIC KWCTIW
S02 COST EFFECTIVENESS (S/TOH)
CATALYST NOx CAPACITY MTtOFIT
LIFE •EDUCTKM FACTO* FACTM
(YtARS) (X) (X) (X)
1.00
t 70 30 1.50
2.00
1.00
1 80 30 1.50
2.00
1.00
t 70 SO 1.50
2.00
1.00
1 80 50 1.50
2.00
1.00
1 70 70 1.50
2.00
1.00
1 80 70 1.50
2.00
1.00
3 70 30 1.50
2.00
1.00
3 80 30 1.50
2.00
1.00
3 70 50 1.50
2.00
1.00
3 80 SO 1.50
2.00
1.00
3 70 70 1.50
2.00
1.00
3 80 70 1.50
2.00
(/TON
«.
(/TON
(/TON
(/TON
(/TON
(/TON
(/TON
(/TON
(/TON
(/TON
(/TON
100
9657.6
10996.2
12334.1
8735.4
10039.9
11344.2
5877.9
6681
7483.7
5320.6
6103.1
6885.8
4263.8
4837.4
S410.8
3862.3
4421.4
4980.4
6058.5
7396.9
8734.9
5586.1
6890.5
8195
3718.5
4521.5
5324.2
3431
4213.6
4996.3
2721.4
3295
3868.4
2512.7
3071.7
3630.8
(OILER XNEUTINO CAPACITY (MU)
300 500 700 1000
8576
9487.1
10398.1
7730.2
8631.8
9533.3
5228
5774.8
6321.4
4716.7
5257.7
5798.6
3799
4189.5
4579.9
3430.3
3816.7
4203.1
4977.1
5888.3
6799.4
4581.2
5482.8
6384.3
3068.7
3615.5
418i>1
2827.3
3368.3
3909.2
2296.7
2647.2
3037.6
2080.8
2467.2
2893.6
8230.3
9055.9
9881.6
7415.9
8236.9
9057.9
5020.4
5515.8
6011.2
4528
5020.6
5513.2
3650.6
4004.4
4358.3
3295.4
3647.2
3999.1
4631.3
5457
6282.6
4266.8
5087.8
5908.8
2861.1
3356.5
3851.9
2638.3
3131.1
3623.7
2108.2
2462
2815.9
1945.8
2297.6
2649.5
8119.6
8908.7
9697.8
7314.1
8100.6
8887.1
4954
5427.4
5900.8
4466.8
4938.7
5410.6
3603.1
3941.2
4279.4
3251.6
1588.7
3925.7
4520.7
5309.8
6098.9
4165.1
4951.6
5738
2794.6
3268.1
3741.5
2577.4
3049.3
3521.2
2060.7
2398.9
2737
1902.1
2239.1
2576.2
8056.6
8818.2
9579.8
7255.2
8015.7
8776.3
4916.1
S373.1
5830
4431.4
4887.8
5344.1
3576
3902.4
4228.8
3226.3
3552.3
3878.2
4457.7
5219.3
5980.9
4106.1
4866.7
5627.2
2756.8
3213.7
3670.7
2542
2998.3
3454.7
2033.6
2360
2686.4
1878.7
2202.7
2528.6
1300
8005.8
8752.7
9499.5
7208.8
7955.4
8702.1
4885.6
S333.8
3781.9
4403.6
4851.5
5299.5
3554.2
3874.3
4194.3
3206.4
3526.4
3846.3
4406.9
S153.8
5900.6
4059.7
4806.4
5553
2726.3
3174.4
3622.3
2514.1
2962.1
3410.1
2011.8
2331.9
2652
1856.8
2176.8
2496.8
G-4
-------
APPENDIX H
Furnace Sorbent Injection
H-l
-------
FUMACt KXKIT INJICTION AND HIMIDIFICATION
Fll*f$f (UUtfif IS»:400 SCA>
CAPii*i cos: (S/U)
fULFM CAJIACITT S02
COMTEK! FACTO! «£t«VAL
<*> (t) (X)
•01LM GEMUATING CAPACITY (Ml)
100 300 500 TOO 1000
1300
1 30
1 50
1 70
2 30
2 50
2 70
3 30
3 50
3 70
4 30
4 50
4 70
30
50
70
30
50
70
30
50
70
30
50
70
30
50
70
30
50
70
30
SO
70
30
50
70
30
30
70
SO
50
70
30
SO
70
50
SO
70
*/kU
*/kU
s/kw
t/kW
*/kU
l/fcu
S/kH
Vku
VKW
t/kv
S/ku
8/W
54
54.9
57.7
54
54.9
57.7
54
54.9
57.7
44.5
48.3
44.3
44.3
45.5
44.4
44.3
45.5
44.4
71.4
73
74.3
71.4
75
74.5
71.4
75
74.5
78.2
79.9
81.4
78.2
79.9
81.4
78.5
79.9
•1.4
34.8
33.2
35.7
34.8
33.2
33.7
54.8
33.3
33.7
41.5
42
42.4
41.5
42
42.4
41.5
42
42.4
41
48.9
49.7
a
48.9
49.7
48.1
48.9
49.7
34.4
59.4
54.3
54.4
53.4
54.5
54.4
39.4
54.5
28.4
28.7
29.1
28.4
28.8
29.1
28.4
28.8
29.1
34.9
39.4
33.9
34.9
39.4
31.9
34.9
35.4
53.9
57
57.7
38.4
57.1
57.8
38.4
57.1
57.8
3*4
41.7
42.4
43.3
41.7
42.4
43.3
41.8
42.4
45.4
24.4
24.9
27.2
24.4
24.9
27.2
24.4
24.9
27.2
52.4
32.9
35.4
32.3
32.9
55.4
32.3
52.9
55.4
54.3
54.9
33.3
54.3
34.9
33.1
54.5
34.9
53.9
58.9
59.7
40.4
59
59.7
40.4
59
59.7
40.4
24.4
24.9
23.1
24.4
24.9
25.1
24.4
24.9
25.1
27.8
28.2
28.4
27.8
28.2
28.4
27.8
28.2
28.4
52.4
52.9
. 55.4
32.4
52.9
33.4
52.4
32.9
55.4
37.3
37.9
584
57.5
38
584
57.5
38
38.4
23.4
23.8
24.1
25.4
23.8
24.1
23.4
23.8
24.1
24.8
27.2
27.4
24.8
27.2
27.4
24.9
27.2
27.4
31.7
32.3
32.7
31.8
32.3
32.7
31.8
52.3
32.8
38.4
39.5
59.9
58.7
59.5
59.9
38.7
39.3
39.9
H-2
-------
WRNACt SORgENT INJECTION AND (MODIFICATION
FSI«€SP (LARGE E»t400 SCA)
ANNUAL COST (•!(!•/UK)
SULFUR CAPACITY S02
CONTENT FACTOR REMOVAL
(X) (X) " <*>
BOILER GENERATING CAP AC ITT (MO
100 300 SOD 700 1000
1100
1
1
1
2
2
2
3
3
3
4
4
4
30
SO
70
30
SO
70
30
SO
70
30
50
70
30
50
70
30
50
70
30
50
70
30
50
70
30
50
70
30
SO
70
30
50
70
30
50
70
30
50
70
30
SO
70
30
50
70
30
50
70
•ids/Mi
•1 lit/Mi
•UU/Mi
•Ult/kVh
•flU/KMi
•UU/Mi
•lUi/kWh
•Ult/Hh
•lilt/Mi
•Ills/Mi
•Uli/kWh
•tUs/Uh
12.1
12.2
12.4
.3
.4
.5
.6
.7
.8
1S.2
15.4
15.6
11
11.2
11.4
9.3
9.4
9.6
18.1
18.4
18.7
13.7
14
14.3
11.9
12.1
12.3
21
21.4
21.8
16.4
16.7
17.1
14.5
14.8
15.1
7.2
7.3
7.4
5.3
5.4
5.5
4.5
4.6
4.7
10.1
10.3
10.5
8
8.2
8.3
7.1
7.2
7.4
13
13.3
13.3
10.7
10.9
11.1
9.7
9.9
10.1
1S.9
16.2
16.6
13.3
13.6
13.9
12.3
12.5
12.8
6
6.1
6.2
4.6
4.7
4.8
4
4.1
4.2
8.9
9.1
9.3
7.3
7.4
7.6
6.6
6.7
6.9
11.4
11.6
11.9
9.7
9.9
10.1
9
9.2
'#4
14.1
14.4
14.7
12.3
12.5
12.8
11.9
11.8
12
5.6
5.7
5.8
4.4
4.4
4.S
3.8
3.9
4
8.5
8.6
8.8
7
7.2
7.3
6.4
6.5
6.7
10.9
11.1
11.3
9.4
9.6
9.8
8.8
9
9.2
13.6
13.9
14.2
12
12.
12.
11.
11.
11.
5.3
5.3
5.4
4.1
4.2
4.3
3.7
3.7
3.8
7.8
8
8.1
.6
.8
.9
.1
.2
.4
10.5
10.7
11
9.2
9.4
9.6
8.
8.
13.
13.
13.
11.7
12
12.3
11.1
11.4
11.6
5.1
5.1
5.2
4
4.1
4.2
3.6
3.7
3.7
7.6
7.8
7-9
6.3
6.6
6.8
6
6.2
6.3
10.4
10.6
10.8
9.1
9.3
9.5
8.5
8.7
8.9
13.3
13.6
13.9
11.8
12.1
12.3
11.1
11.4
11.7
H-3
-------
FUUUCt SOUU1 INJECTION AND MUMMIFICATION
FS!*€$>> (uutoe E»:4oo sou
$02 COST EFrecrivewn <«/TON>
SULFUR CAPACITY $02
CONTENT ttCTOt . «ENOVAl
(X) (X) (*)
1
1
1
2
2
2
3
5
3
4
'• 4
4
30
SO
70
30
SO
70
30
SO
70
30-
SO
70
30
SO
70
30
SO
70
30
SO
70
30
SO
70
30
SO
70
30
SO
70
30
SO
70
30
SO
70
30
SO
70
30
SO
70
30
SO
70
30
SO
70
I/TOI
I/TON
•/TON
I/TOH
•/TON
.(/TOM
S/TON
»/TCB
•/TON
S/TOH
•/TON
S/TOi
80ILEI CCNEIATING CAPACITY
-------
APPENDIX I
Natural Gas Reburn
1-1
-------
NATUUL GAI UKJM
CAPITAL COST (S/tW>
FUEL MICE CAPACITY fUtl
0IFFCRENTIAL fACTM SUMTITUTCD
<»> . (X)
1 S/IMTU
1 S/IMTU
i S/IMTU
2 I/IMTU
2 S/IMTU
2 S/IMTU
3 S/IMTU
3 S/IMTU
3 i/wwrv
30 IS
a
3
SO 1S
29
70 IS
29
30 IS
29
S
SO IS
29
S
70 IS
29
S
30 IS
29
S
SO IS
29
S
70 IS
29 •
MILCH CEttRATIM CAPACITY *
S/ku 23.1
27.9
18.3
S/kU 23.1
27.9
18.3
S/kU 23.1
27.9
19.4
S/kH 26.9
34.2
19.5
S/kW 26.9
34.2
19.S
»/kW 26.9
S4.2
1S.7
16
11.4
13.7
16
13.7
16
19 A
I*»O
17.4
22.2
12.6
17.4
22.2
12.6
17.4
22.2
13.9
21.2
28.5
13.9
21.2
28.5
13.9
21.2
28.5
11.8
14.1
9.3
11.8
14.1
11.8
14.1
15. S
20.1
10.7
15.5
20.3
10.7
15.5
20.3
12
19.3
26.6
12
19.3
26.6
12
19.3
3i.6
10.7
13
8.4
10.7
13
10.7
13
14.S
19.3
9.7
U.S
19.3
9.7
U.S
19.3
10.9
18.2
29.3
10.9
18.2
29.5
10.9
18.2
29.9
9.8
12.1
7.5
9.8
12.1
9.8
12.1
13.5
18.3
8.7
13.5
18.3
8.7
13.9
18.3
10
17.3
24.6
10
17.3
24.6
10
17.3
24.6
1300
9.1
11.4
6.8
9.1
11.4
9.1
11.4
12.9
17.7
8.1
12.9
17.7
8.1
12.9
17.7
9.3
16.
23.
9.
16.
23.
9.
16.
23.
1-2
-------
NATURAL GAS RCUU
ANNUAL COST (•UU/kttl)
FUEL PRICE CAPACITY FUEL
DIFFERENTIAL FACTOR SUBSTITUTED
. (X)
1 */wenj
1 S/MSTU
1 S/MQTU
2 S/WUTU
2 S/WttTU
2 S/MttTU
3 S/IW8TU
3 S/IMTU
3 S/i««TU
5
30 15
25
5
50 15
25
5
70 15
25
5
30 15
23
5
50 15
25
5
70 15
23
5
30 15
25
5
50 15
25
5
70 15
25
BOILER GENERATING CAPACITY (MO
100 300 500 700 1000
1.
•< ill/Ml 3.
5.
1.
•llU/Mft 3.
5-*
1.4
•UK/Ml 3.3
5.2
3
•ItU/Wh 7.2
11.3
2.6
•ills/Mi 6.6
10.6
2.4
•UU/kWh 6.4
10.3
4.1
•UU/Mi 10.4
16.7
3.6
•UU/Uh 9.7
15.9
3.4
•tlls/bft 9.3
15.5
1.5
3.5
5.5
1.3
3.2
3.1
1.2
3.1
3
2.6
6.7
10.9
2.3
6.3
10.4
2.2
6.2
10.1
3.6
10
16.3
3.4
9.
13.
3.
9.
13.
1.3
3.3
5.3
1.2
3.1
5
1.1
3
4.9
2.4
6.6
10.7
2.2
6.2
10.3
2.1
6.1
10.1
3.5
9.8
16.1
3.3
9.4
13.5
3.2
9.2
15.2
1.2
3.2
5.2
1.1
3
5
1.1
3
4.9
2.3
6.3
10.6
2.2
6.2
10.2
2.1
6.1
10
3.4
9.7
16
3.2
9.3
13.4
3.1
9.2
15.2
1.2
3.2
5.1
1.1
3
4.9
1
2.9
4.8
2.2
6.4
10.3
2.1
6.1
10.2
2.1
6
10
3.3
9.6
15.9
3.2
9.3
15.4
3.1
9.1
13.2
1300
1.1
3.1
5.1
1
3
4.9
1
2.9
4.8
2.2
6.3
10.5
2.1
6.1
10.1
2
6
10
3.3
9.6
15.9
3.1
9.2
15.4
3.1
9.1
13.1
1-3
-------
BATUMI CM UKM
WK COST EFFECTIVOMI (t/T(M)
nti pticc CAPACITY net
OirrUCMTIAL PACTOt SUBSTITUTED
(X) («)
1 S/IMTU
1 S/IMTU
1 S/IMTU
2 S/IMTU
2 S/IMTU
2 S/IMTU
3 S/IMTU
3 S/IMTU
3 «/>wrj
5
30 15 S/TOH
25
5
50 11 S/TOH
25
5
70 19 S/TOH
25
5
30 15 S/TOH
25
5
50 15 S/TOH
25
S
70 15 S/TOH
25
5
30 15 S/TW
25
5
50 15 S/TOH
25
5
70 15 S/TW
25
80ILO GEJtEIATIK CAPACITY
100 300 500 700
747.
1510.
2274.
sav.
1329.
2069.
521.
1251.
itei.
1161.
2754.
4346.
991.
2534.
4077.
570
1333.8
2097.6
483
1223.3
1963.5
445.7
1175.9
1906
984.4
2577.1
4169.5
884.6
2428
3971.4
917.9 841.8
2440.1 2364.1
3962.4 3886.5
1576.4 1398.8
3997.2 3820.2
6418.8 6241.3
1392.7 1286.2
3739.1 3632.7
6085.8 5979.3
1314 1237.9
3628.4 3352.4
5942.7 5866.9
510.7
1274.3
2038.4
447.4
1187.7
1927.9
420.2
1150.4
1880.6
925
2517.7
4110.3
849
2392.4
3935.9
816.3
Z338.7
3861
1339.4
3760.8
6182.3
1250.6
3597.1
5943.7
1212.4
3327
5844.5
477.8
1241.6
2005.4
427.6
1167.9
1908.1
406.2
1136.3
1866.5
892.1
2404.8
4077.4
829.2
2372.6
3916.1
802.2
2324.6
3846.9
1306.5
3727.9
6149.3
1230.8
3577.4
5924
1198.3
3512.9
5827.4
(NO
1000
447.4
1211.2
1975.1
409.4
1149.6
1890
393.1
1123.3
1853.5
861.7
2454.4
4047
811
2354.4
3697.9
789.2
2311.6
3833.9
1276.1
3697.5
6119
1212.5
3559.2
5905.8
1185.3
3499.8
5814.4
1300
427.
1191.
1953.
397.
1137.
1878.1
384.7
1114.8
1845
842
2434.6
4027.3
799.1
2342.5
3886
780.7
2303.1
3825.5
1256.4
3677.8
6099.2
1200.7
3547.3
5893.9
1176.8
3491.4
5805.9
1-4
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