EPA/600/R-00/093
November 2000
CONTROLLING SO2 EMISSIONS:
A REVIEW OF TECHNOLOGIES
Prepared by:
Ravi K. Srivastava
U.S. Environmental Protection Agency
National Risk Management Research Laboratory
Research Triangle Park, NC 27711
Prepared for:
U.S. Environmental Protection Agency
Office of Research and Development
Washington, B.C. 20460
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Notice
This report has been peer and administratively reviewed by the U.S. Environmental Protection
Agency, and approved for publication. Mention of trade names or commercial products does not
constitute endorsement or recommendation for use.
This document is available to the public through the National Technical Information Service,
Springfield, Virginia 22161.
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Foreword
The U.S. Environmental Protection Agency is charged by Congress with protecting the Nation's
land, air and water resources. Under mandate of national environmental laws, the Agency strives
to formulate and implement actions leading to a compatible balance between human activities and
the ability of natural systems to support and nurture life. To meet this mandate, EPA's research
program is providing data and technical support for solving environmental problems today and
building a science knowledge base necessary to manage our ecological resources wisely,
understand how pollutants affect our health, and prevent or reduce environmental risks in the
future.
The National Risk Management Research Laboratory is the Agency's center for investigation of
technological and management approaches for reducing risks from threats to human health and the
environment. The focus of the Laboratory's research program is on methods for the prevention and
control of pollution to air, land, water and subsurface resources; protection of water quality in
public water systems; remediation of contaminated sites and ground water; and prevention and
control of indoor air pollution. The goal of this research effort is to catalyze development and
implementation of innovative, cost-effective environmental technologies; develop scientific and
engineering information needed by EPA to support regulatory and policy decisions; and provide
technical support and information transfer to ensure effective implementation of environmental
regulations and strategies.
This publication has been produced as part of the Laboratory's strategic long-term research plan. It
is published and made available by EPA's Office of Research and Development to assist the user
community and to link researchers with their clients.
E. Timothy Oppelt, Director
National Risk Management Research Laboratory
in
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Abstract
Sulfur dioxide (862) scrubbers may be used by electricity generating units to meet the
requirements of Phase II of the Acid Rain 862 Reduction Program. Additionally, the use of
scrubbers can result in reduction of mercury and particulate matter emissions. It is timely,
therefore, to review commercially available flue gas desulfurization (FGD) technologies that have
an established record of performance.
The review of FGD technologies presented in this report describes these technologies,
assesses their applications, and characterizes their performance. Additionally, the report describes
some of the advances that have occurred in FGD technologies. Finally, the report presents an
analysis of the costs associated with applications of limestone forced oxidation, lime spray dryer,
and magnesium-enhanced lime FGD processes. The information presented in this paper should be
useful to parties evaluating FGD technology applications.
Acknowledgements
We acknowledge the invaluable contributions of Wojciech Jozewicz and
Carl Singer under EPA Contract 68-C-99-201 with ARCADIS Geraghty
& Miller, Inc., P.O. Box 13109, Research Triangle Park, NC 27709.
IV
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Contents
Abstract iv
Acknowledgements iv
List of Figures viii
List of Tables ix
List of Symbols x
List of Acronyms and Abbreviations xii
Conversion Table - English Units to SI Units xiii
Chapter 1 Introduction 1
Chapter 2 FGD Technology 3
Introduction 3
Wet FGD Technologies 4
Limestone Forced Oxidation 7
Limestone Inhibited Oxidation 10
Lime and Magnesium-Lime 10
Seawater Process 11
Dry FGD Technologies 11
Lime Spray Drying 12
Duct Sorbent Inj ection 14
Furnace Sorbent Inj ection 14
Circulating Fluidized Bed 17
Regenerable FGD Technologies 17
Wet Regenerable FGD 17
Sodium Sulfite 17
Magnesium Oxide 17
Sodium Carbonate 19
Amine 19
Dry Regenerable FGD 19
Activated Carbon 19
Chapter 3 Technology Applications 20
Introduction 20
Historical Applications 20
Current Application 24
Chapter 4 Performance 32
Introduction 32
SO2 Removal Efficiency 32
Energy Requirements 33
Applicability 35
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Chapter 5 Advances 36
Introduction 36
Once-through Wet FGD Technology 36
Ammonia Scrubbing 39
Chapter 6 FGD Cost 43
General Approach 43
Limestone Forced Oxidation 43
Capital Cost 46
Reagent Feed Area 46
SO2 Removal Area 48
Flue Gas Handling Area 50
Waste/By-product Handling Area 51
Support Equipment Area 52
Total Capital Requirement 52
Operation and Maintenance Cost 54
Validation 55
State-of-the-art Model 57
Lime Spray Drying 64
Sensitivity Analysis 64
Capital Cost 64
Reagent Feed Area 64
SO2 Removal Area 66
Flue Gas Handling Area 67
Waste/By-product Handling Area 68
Support Equipment Area 68
Total Capital Requirement 69
Operation and Maintenance Cost 69
Validation 70
State-of-the-art Model 71
Magnesium-enhanced Lime 77
General Approach 77
Capital Cost 77
Reagent Feed Area 77
SO2 Removal Area 77
Flue Gas Handling Area 79
Waste/By-product Handling Area 80
Support Equipment Area 81
Total Capital Requirement 81
Operation and Maintenance Cost 81
State-of-the-art Model 82
Summary of FGD Cost 84
Chapter 7 Additional Benefits 88
VI
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Introduction 88
Once-through Wet FGD 89
DryFGD 90
References 92
Vll
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List of Figures
2-1. FGD technology tree 4
2-2. Baseline wet FGD system 5
2-3. Lime spray dryer FGD system 13
2-4. Schematic of DSI 15
2-5. Schematic of FSI 16
2-6. Schematic of CFB 18
3-1. Historical application of FGD technology in the United States 21
3-2. Historical application of FGD technology throughout the world 22
3-3. Wet FGD technology application in the United States 23
3-4. Dry FGD technology application in the United States 25
3-5. Regenerable FGD technology application in the United States 26
3-6. Percent shares (capacity) of the three FGD technologies installed 27
3-7. Comparison of limestone and non-limestone wet FGD applications 30
4-1. Design SC>2 removal efficiencies for wet limestone and spray drying processes 34
6-1. Schematics of LSFO system's equipment areas 47
6-2. Comparison of model predictions with cost data for LSFO 56
6-3. Comparison of LSFO Cost Model to IPM model predictions for 2 to 4 percent
sulfur coal 58
6-4. TCR predictions for 2 to 4 percent sulfur coal by LSFO SUSCM 60
6-5. Fixed O&M predictions for 2 to 4 percent sulfur coal by LSFO SUSCM 62
6-6 Variable O&M predictions for 2 to 4 percent sulfur coal by LSFO SUSCM 63
6-7. Validation ofLSD cost model 72
6-8 LSD TCR predictions by LSD SUSCM 74
6-9. LSD fixed O&M predictions by LSD SUSCM 75
6-10. LSD variable O&M predictions by LSD SUSCM 76
6-11. MEL TCR predictions by MEL SUSCM 85
6-12. MEL fixed O&M predictions by MEL SUSCM 86
6-13. MEL variable O&M predictions by MEL SUSCM 87
Vlll
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List of Tables
3-1. Coal-fired Electrical Generation Capacity (MWe) Equipped with FGD
Technology (1998) 28
3-2. Total Capacity (MWe) Equipped with Wet FGD Technology (1998) 29
3-3. Total Capacity (MWe) Equipped with Dry FGD Technology (1998) 29
3-4. Number of Installed FGD Technology Systems (1998) 31
3-5. Average Size (MWe) of FGD Technology Systems (1998) 31
4-1. Design SC>2 Removal Efficiencies 33
5-1. Advanced Options for New Wet FGD Scrubbers 37
6-1. Sensitivity Analysis of LSFO Annual Operating Cost (baseline cost of
10.31 mills/kWh) 45
6-2. Representative Values for LSFO Variables with Minor Cost Impacts 46
6-3. TCR Calculation Method 53
6-4. Financial Factors for FGD Construction, Constant Dollars 54
6-5. Model Validation Summary for LSFO FGD (1994 Dollars) 56
6-6. "State-of-the-art" LSFO Design Decisions 59
6-7. Sensitivity Analysis of LSD Annual Operating Cost (baseline value of
10.02 mills/kWh) 65
6-8. Representative Values for Variables with Minor Cost Impacts 66
6-9. Validation of LSD Model 70
6-10. "State-of-the-art" LSD Design Decisions 73
6-11. "State-of-the-art" MEL Design Decisions 83
6-12. Cost in 1998 Constant Dollars for Selected FGD Technologies 84
IX
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List of Symbols
Symbol
A&S
ABSORBER
ABSORBER 1
ABSORBER 2
ACFM
ACFM1
ACFM2
ACFM3
BARE MODULEE
BARE MODULEo
BARE MODULER
BARE MODULEw
BM
BME
BMF
BMG
BMR
BMW
CB&H
CcaC03
CcaO
CDBA
CDL
CDS
CF
CHIMNEY
CREDIT
Meaning
Administration and support cost
Absorber cost
RLCS absorber cost
Alloy absorber cost
Flue gas flow into absorber
Flue gas flow out of absorber
Flue gas flow out of ID fans
Flue gas flow out of particulate control device
Support equipment area auxiliary cost
Flue gas handling area auxiliary cost
SO2 removal area auxiliary cost
Waste handling area auxiliary cost
Capital cost for FGD system
Capital cost component for support equipment area
Capital cost component for reagent feed area
Capital cost component for waste handling area
Capital cost component for waste handling area
Capital cost component for waste handling area
Cost of ball mill and hydroclones
Cost of limestone
Cost of lime
Cost of DBA tank
Cost of disposal with landfilling
Cost of disposal with gypsum stacking
Capacity factor
Cost of chimney
By-product credit
Unit
dollars
dollars
dollars
dollars
cfm
cfm
cfm
cfm
dollars
dollars
dollars
dollars
dollars
dollars
dollars
dollars
dollars
dollars
dollars
dollars
dollars
dollars
dollars
dollars
%
dollars
Dollars
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Symbol
FAFDC
Fd
FGPM
Fixedo&M
FRL
FRso2
FTCE
HHV
HR
ID FANS
L/G
ML&M
Na
Nf
NP
OL
P
POWER
PUMP
PUMPS
SPRAY DRYERS
SPRAY DRYERS 1
SPRAY DRYERS2
STEAM
TER
THICKENER
Variableo&M
Wt%S
Meaning
Allowance for funds during construction factor
F-factor
Slurry flow rate
Fixed operation and maintenance cost
Reagent feed rate
SO2 feed rate to the FDG System
Total cash expended factor
Coal heating value
Plant heat rate
Cost of ID fans
Liquid-to-gas ratio
Maintenance, labor, and materials cost
Number of absorbers
Number of fans
Number of pumps
Operating labor cost
Percent oxygen in the stack
Cost of electrical energy
Cost of pump
Cost of pumps
Cost of spray dryers
Cost of RLCS spray dryers
Cost of alloy spray dryers
Cost of steam
Thermal energy required
Cost of thickener
Variable operation and maintenance cost
Coal sulfur content
Unit
scf/106 Btu
gal/min
dollars
Ib/hr
Ib/hr
Btu/lb
Btu/kWh
dollars
gal/1000 ft3
dollars
dollars
%
dollars
dollars
dollars
dollars
dollars
dollars
dollars
Btu
dollars
dollars
Wt%
XI
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List of Acronyms and Abbreviations
CAAA
CFB
CUECost
DBA
DSI
EPA
ESP
FGD
FSI
ID
L/G
LSD
LSFO
LSIO
MEL
MWe
NAAQS
O&M
PM2.5
RLCS
SUSCM
TCR
TPC
TPI
WESP
Clean Air Act Amendments of 1990
Circulating Fluidized Bed
Coal Utility Environmental Cost Workbook
Dibasic Acid
Duct Sorbent Injection Process
United States Environmental Protection Agency
Electrostatic Precipitator
Flue Gas Desulfurization
Furnace Sorbent Injection Process
Inside Diameter
Liquid-to-gas Ratio
Lime Spray Drying Process
Limestone Forced Oxidation Process
Limestone Inhibited Oxidation Process
Magnesium-Enhanced Lime Process
Unit Electrical Generating Capacity
National Ambient Air Quality Standards
Operation and Maintenance
Particulate Matter Less than 2.5 • m (Aerodynamic Diameter)
Rubber-Lined Carbon Steel
State-of-the-art Utility Scrubber Cost Model
Total Capital Requirement
Total Plant Cost
Total Plant Investment
Wet Electrostatic Precipitator
Xll
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To Obtain
m
m2
m3
°C
kg
J/kg
m3/s
m3/s
J/kWh
mills
2
kg/nT
Conversion Table - English Units to SI Units
From Multiply by
ft 0.3048
ft2 9.29 • 10'2
ft3 2.83 • 10'2
°F 5/9 (°F - 32)
Ib 0.454
Btu/lb 1.33 • 10
cfm 4.72 • 10
gpm 6.31 • 10
Btu/kWh 1055.056
$ 0.001
in. Hg 345.31
-4
-4
-5
Xlll
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CHAPTER 1
INTRODUCTION
Combustion of sulfur-containing fuels, such
as coal and oil, results in sulfur dioxide (862)
formation. SC>2 emissions are known to cause
detrimental impacts on human health and the
environment. The major health concerns
associated with exposure to high
concentrations of SC>2 include breathing
difficulties, respiratory illness, and
aggravation of existing cardiovascular
disease. In addition to the health impacts,
SC>2 leads to acid deposition in the
environment. This deposition causes
acidification of lakes and streams and damage
to tree foliage and agricultural crops.
Furthermore, acid deposition accelerates the
decay of buildings and monuments. While
airborne, 862 and its particulate matter
derivatives contribute visibility degradation.
Electric power generating units account for
the majority of SC>2 emissions in the U.S. In
1998, these units contributed 64 percent of
the national SC>2 emissions.1 To mitigate SC>2
emissions from electric power generating
units, the Acid Rain 862 Reduction Program2
was established under Title IV of the Clean
Air Act Amendments of 1990 (CAAA). This
two-phase program was designed to reduce
SC>2 emissions from the power generating
industry.
Phase I of the Acid Rain SC>2 Reduction
Program began on January 1, 1995, and
ended December 31, 1999. In 1997, 423
power generating units, affected under Phase
I, emitted 5.4 million tons of SC>2 (1.7 million
tons below the allowable 7.1 million tons of
SO2).3 Thus, the SO2 emissions in 1997
reflect an output of 23 percent below the
allowable amount.
Phase II of the Acid Rain 862 Reduction
Program began on January 1, 2000. The
nationwide cap for SC>2 will be 9.48 million
tons from 2000 through 2009. In 2010, the
cap will be reduced further to 8.95 million
tons, a level approximately one-half of
industry-wide emissions in 1980. To meet
the requirements of this phase, some power
generating units may use FGD technologies.
Additionally, the use of these technologies
can result in the reduction of fine particle
precursor emissions and mercury emissions
from combustion units. It is timely, therefore,
to examine the current status of FGD (or 862
scrubbing) technologies.
This report presents a review of current FGD
technologies. Following the introduction,
Chapter 2 presents a concise review of
commercially available FGD technologies.
Technology applications on combustion units
in the United States and abroad are discussed
in Chapter 3. The performance and
applicability of the most commonly occurring
types of FGD technology installations is
presented in Chapter 4. A review of recently
reported technical advances to FGD
technologies is provided in Chapter 5.
Capital and operating costs of LSFO, LSD,
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and MEL are analyzed in Chapter 6.
Additional benefits achieved with wet
limestone scrubbers and spray dryers are
discussed in Chapter 7. References reviewed
and utilized for the production of this report
are given at the end.
It is expected that this review will be useful to
a broad audience, including: (1) individuals
responsible for developing and implementing
SC>2 control strategies at sources, (2) persons
involved in developing SCh and other
regulations, (3) State regulatory authorities
implementing SCh control programs, and (4)
interested public at large. Moreover, persons
engaged in research and development efforts
aimed at improving cost-effectiveness of
FGD technology may also benefit from this
review.
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CHAPTER 2
FGD TECHNOLOGY
Introduction
Various technologies exist that have been
designed to remove 862 from flue gas
produced by electricity generating plants.
These technologies represent a varying
degree of commercial readiness. Some can
claim tens of thousand of hours of operational
experience, while others have only recently
been demonstrated at commercial plants.
This report considers only commercially
available FGD technologies that have an
established record of reliable performance
and sufficient quality and quantity of data to
determine the cost of their deployment.
Commercially available FGD technologies
can "conventionally" be classified as once-
through and regenerable, depending on how
sorbent is treated after it has sorbed SC>2.4 In
once-through technologies, the SC>2 is
permanently bound by the sorbent, which
must be disposed of as a waste or utilized as a
by-product (e.g., gypsum). In regenerable
technologies, the SC>2 is released from the
sorbent during the regeneration step and may
be further processed to yield sulfuric acid,
elemental sulfur, or liquid 862. The
regenerated sorbent is recycled in the SC>2
scrubbing step. Both once-through and
regenerable technologies can be further
classified as wet or dry. In wet processes, wet
slurry waste or by-product is produced, and
flue gas leaving the absorber is saturated with
moisture. In dry processes, dry waste
material is produced and flue gas leaving the
absorber is not saturated with moisture.
Depending on process configuration and local
market conditions at the plant site, once-
through wet FGD processes can produce
slurry waste or salable by-product. This
waste/by-product must be dewatered in some
fashion prior to disposal or sale (in case of a
salable by-product). The "conventional"
classification of FGD processes is shown in
Figure 2-1.
A review of FGD technology applications
was conducted based on the information
provided in CoalPowerS Database, available
from the International Energy Agency's Coal
Research Centre in London, England. This
database lists commercial FGD applications.
The review reveals that regenerable FGD
processes are being used only marginally,
with once-through FGD processes involved in
the vast maj ority of applications. Therefore,
for this work, FGD technologies were
grouped into the following three major
categories:
• Wet FGD (composed of once-through wet
FGD)
• Dry FGD (composed of once-through dry
FGD)
• Regenerable FGD (composed of wet and
dry regenerable FGD)
The above grouping of FGD technologies is
consistent with other evaluations of FGD,5
and will be used in the remaining chapters of
this report. Accordingly, when wet FGD is
mentioned in the remainder of this report, it is
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1
Wet
Flue Gas Desulfurization
i
i
Once-through
I
1
Regenerable
I
I I I
— Limestone Forced Oxidation
— Limestone Inhibited Oxidation
— Lime
— Magnesium-Enhanced Lime
— Seawater
Dry
— Lime Spray Drying
— Duct Sorbent Injection
— Furnace Sorbent Injection
— Circulating Fluidized Bed
Wet Dry
— Sodium Sulfite 1 — Activated Carbon
— Magnesium Oxide
— Sodium Carbonate
— Amine
Figure 2-1. FGD technology tree.
meant as once-through wet FGD. Similarly,
when dry FGD is mentioned, it is meant as
once-through dry FGD. Moreover, as
regenerable technologies are used only
marginally, their coverage in this report is
limited.
Wet FGD Technologies
In wet FGD processes flue gas contacts
alkaline slurry in an absorber. The absorber
may take various forms (spray tower or tray
tower), depending on the manufacturer and
desired process configuration. However, the
most often-used absorber application is the
counter-flow vertically oriented spray tower.
A diverse group of wet FGD processes have
evolved to take advantage of particular
properties of various sorbents and/or by-
products. All wet FGD processes discussed
here are once-through (i.e., non-regenerable).
A generalized flow diagram of a baseline wet
FGD system is shown in Figure 2-2.
SCVcontaining flue gas is contacted with
limestone slurry in an absorber. Limestone
slurry is prepared in two consecutive steps.
First, limestone is crushed into a fine powder
with a desired particle size distribution. This
takes place in a crushing station; e.g., ball
mill (fine crushing maximizes the dissolution
rate of a given limestone). Next, this fine
powder is mixed with water in a slurry
preparation tank. Sorbent slurry from this
tank is then pumped into the absorber
reaction tank.
As mentioned before, the absorber is most
often a counterflow tower with flue gas
flowing upwards, while limestone slurry is
sprayed downwards by an array of spray
nozzles. In the absorber, SC>2 is removed by
both sorption and reaction with the slurry.
Reactions initiated in the absorber are
completed in a reaction tank, which provides
retention time for finely ground limestone
particles to dissolve and to react with the
dissolved 862.
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FGD TECHNOLOGY
Limestone
Water
Crushing
Station
Slurry
Preparation
Tank
Rue Gas
In
_r \
Rue Gas
Out
Absorber
Chimney
Process Water
Disposal
Dewatering
Figure 2-2. Baseline wet FGD system.
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The overall reactions in the absorber and in
the reaction tank can be summarized by:
SO2+CaCO3+y2H2O
CaSO3»/2H2
and
SO2 +/2O2+ CaCO3 + 2H2O
CaSO4 • 2H2O + CO2
(2-1)
(2-2)
The complex chemistry summarized by the
above equations involves SO2-CO2-H2O
equilibrium relationships in the absorber,
limestone dissolution, and sulfite/sulfate
crystallization (occurring mostly in the
reaction tank)6. If the oxidation of sulfite to
sulfate is not controlled, the wet limestone
system is operating under the so-called
natural oxidation. Depending on SC>2
concentration and the excess air in the flue
gas, as well as on slurry pH, some systems
may be operated in the natural oxidation
mode. However, for most applications, it is
beneficial to control oxidation.
The dissolution and crystallization reactions
in the reaction tank are, to a large extent,
controlled by the pH of the liquid, which is a
function of limestone stoichiometry (number
of mols of Ca added per mol of SC>2
removed). Both pH and limestone
stoichiometry are preset parameters for the
operation of an absorber. Normally, the
required stoichiometry of a limestone wet
FGD system varies from 1.01 to 1.1 moles of
CaCO3 per mole of SO2 (1.01 to 1.05 for
modern scrubbers) and pH is in the range 5.0
to 6.0. A gradual decrease in a preset
operating value of pH indicates increased
limestone consumption and triggers the fresh
limestone feed. Spent sorbent from the
reaction tank (slurry bleed) is dewatered and
disposed of in a waste slurry pond (ponding).
The complexity of the dewatering process is
determined by the chemical composition and
crystal habit of the spent sorbent, and whether
the end product is to be utilized or
discharged. For example, CaSC>4 is easier to
dewater than
Entrained slurry droplets that escaped from
the absorber's spray zone and were carried
out by the flue gas are separated in an
impaction-type mist eliminator. Mist
eliminator design parameters include style
(chevrons, mesh pads, baffles, etc.), blade
number and spacing, and wash system
configuration. The mist eliminator plays an
important role in preventing corrosion of
downstream equipment and ducts, as well as
deposition of stack effluent in the immediate
vicinity of the plant. Mist eliminators can be
designed for either a vertical or horizontal
configuration. A horizontal configuration
offers several advantages over a vertical
configuration; e.g., better drainage.
However, the drawbacks of horizontal mist
eliminators include increased flue gas
pressure drop and more difficult maintenance.
Wet FGD process variables include: flue gas
flow rate, liquid-to-gas ratio (L/G), recycle
slurry pH, flue gas SC>2 concentration, and
solids concentration and retention time. The
effect of these variables on the operation of a
wet FGD system is discussed below.
Flue gas velocity optimization considerations
depend on the type of wet absorber used.
Normally, the upper limit for flue gas
velocity in a counterflow absorber depends
on the capability of the mist eliminators to
prevent droplet carryover.7 Droplet
carryover, or droplets escaping from the unit
eliminator, can increase duct corrosion
downstream of the absorber. Some absorbers
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have a perforated tray added for the
improvement of SC>2 capture. In such cases,
the optimum flue gas velocity is determined
by the tray design. For this type of absorber,
excessive flue gas velocity will cause an
absorber to "flood," whereas too low a
velocity will prevent slurry holdup on the
tray. For a given scrubber, trays are designed
for a maximum gas velocity, so as not to
flood.
Another type of wet FGD absorber that could
be used for 862 control is a packed absorber.
Packed absorber utilizes a material placed in
it to provide a surface over which scrubbing
solution is distributed. In this manner,
gas/liquid contact surface area is generated.
As far as a mist eliminator's operation is
concerned, higher flue gas velocities could be
used for a packed absorber without causing
its failure and a subsequent droplet carryover.
Packed absorbers can be used only for clear
solution systems (systems with a scrubbing
medium being a solution rather than a slurry).
L/G is usually expressed in terms of gallons
of slurry per 1000 ft3 of flue gas at actual
conditions leaving the absorber. The amount
of surface system available for the reaction
with SC>2 is determined by L/G. For a
counterflow spray absorber operated at a
given flue gas flow rate, L/G approximates
the surface area of droplets and is one of the
main design variables available to obtain a
desired SC>2 removal in the absorber. The
amount of available alkalinity for the reaction
with SC>2 increases with the increasing L/G.
L/G also affects the oxidation rate of
sulfite/bisulfite reaction products in the
absorber by affecting the absorption rate of
62 from the flue gas. As will be explained
further in this report, oxidation rate affects
the potential for scaling absorber internals.
Slurry pH also has a significant effect on 862
removal efficiency in a wet FGD system. In
addition, pH is likely the single most
important control variable for absorber
operation. It determines the amount of
limestone added to the system. Within the
operational range, increasing the amount of
limestone added increases the amount of 862
removal. This is because of the increased
concentration of soluble alkaline species and
undissolved reagent. This reagent is then
available for dissolution and renewal of
alkalinity in the liquid phase.
At constant operating conditions of a
scrubber, increasing the concentration of 862
(increasing sulfur content of fuel) will
decrease SC>2 removal efficiency by a wet
absorber. This decreased efficiency is
observed because increasing 862
concentration causes a more rapid depletion
of liquid phase alkalinity causing the increase
of liquid phase resistance.
Solids concentration and retention time affect
the reliability of wet FGD operation. Solids
concentration in the slurry is typically
maintained at 10 to 15 percent solids by
weight. It is controlled by removing a part of
the slurry from the reaction tank for
subsequent dewatering. Proper solids
concentration in the slurry is necessary to
ensure scale-free operation of the absorber.
Correct solids retention time in the reaction
tank is essential to achieving high utilization
of limestone and maintaining correct handling
and dewatering properties of solids. Typical
solids retention time for wet FGD is 12 to 14
hours.7
Limestone Forced Oxidation
As described above, wet FGD can be
operated reliably in a natural oxidation mode
under certain favorable conditions. However,
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for the majority of applications, it is
necessary to control the extent of oxidation in
order to improve operational reliability of the
system. Over the years, several process
variations have been designed to improve the
operational reliability of wetFGD
technology. Consequently, the limestone
forced oxidation process (LSFO) has become
the preferred FGD technology worldwide.
First-generation wet limestone FGD systems
were plagued with scaling problems, resulting
from oxidation of the reaction products to
calcium sulfate (gypsum) that would deposit
throughout the absorber, mist eliminator, and
piping. Gypsum scale typically forms via
natural oxidation when the fraction of
calcium sulfate in the slurry (slurry oxidation
level) is greater than 15 percent. Initially,
gypsum scaling was combated by installation
of extra capacity.
One way to prevent the scaling problem is to
blow air into the absorbent slurry to
encourage controlled oxidation outside of the
absorber. This type of FGD system,
limestone forced oxidation, provides rapid
calcium sulfate crystal growth on seed
crystals. LSFO minimizes scaling in the
scrubber and also results in slurry that can be
more easily dewatered. Consequently, the
LSFO system has become the preferred
technology worldwide. The most often used
configuration is for the air to be blown into
the reaction tank (in-situ oxidation).
Alternatively, air can be blown into an
additional hold tank (ex-situ oxidation).
LSFO requires compressors/ blowers and
additional piping, compared to a system
without forced oxidation.
The prime benefit of scale control derived
from forced oxidation is greater scrubber
absorber availability. As a result, the need
for redundant capacity is greatly reduced.
The added benefits are the formation of a
stable product, a salable by-product (which
eliminates the need for landfilling), and
smaller dewatering equipment. Nearly
complete (99 percent plus) oxidation is
required for a commercial quality by-product.
This level of oxidation can be accomplished
in a modern wet FGD system. However, the
salability of the wet FGD by-product (FGD
gypsum) is also a function of the demand for
gypsum. Depending on site-specific
conditions, LSFO may produce a salable by-
product in the form of commercial grade
gypsum that could be used for wallboard
manufacturing. When salable gypsum is not
attainable, dry FGD waste is piled (gypsum
stacking) or landfilled. Gypsum stacking is
the procedure where a gypsum slurry is sent
to the stacking area, allowed to have the
solids to separate from the water, and then
removing the water and leaving the solids as
a pile.
The solids handling system for LSFO consists
of primary and secondary dewatering, solids
modification unit, and ultimate waste
disposal, regardless whether a part or all of
the by-product will be sold as commercial
quality gypsum. The objective of primary
dewatering is to increase the solids
concentration of spent limestone slurry from
the reaction tank discharge conditions (10 to
15 percent by weight) to between 30 and 50
percent by weight. Primary dewatering is
accomplished by hydroclones. The process
water recovered during primary dewatering is
recirculated to the absorber. Solids
discharged from the primary dewatering unit
are directed to the underflow storage tank.
The objective of secondary dewatering is to
reduce the moisture content (increase solids
-------
content) beyond the setpoint of primary
dewatering. The solids content of the
material leaving this stage will be 45 to 90
percent. This relatively wide range of solids
concentration in the product of secondary
dewatering is a result of different disposal
methods for the product. For an LSFO
absorber aimed at commercial gypsum
production, solids concentration in the
product will be in the high end of the range.
However, for an absorber operated as LSFO,
but without product recovery, the solids
concentration will be at the low end of the
range.
The types of equipment most often used for
secondary dewatering are belt and/or drum
vacuum filters and centrifuges. The selection
of the equipment depends on the quality of
product desired. If commercial quality
gypsum is desired, then belt vacuum filters
may be selected over drum filters because of
their ability to provide superior cake washing
capabilities (important to achieve gypsum
specifications). The process water recovered
during secondary dewatering is recirculated
to the absorber.
Solids discharged from the secondary
dewatering unit are directed either to the
modification unit of solids handling or to the
temporary storage system. During the
modification, solids are stabilized or fixated
to improve their strength bearing, landfill,
and leachate characteristics. This is most
often accomplished by mixing dewatered
solids with fly ash and lime in a pug mill to
promote the pozzolanic reaction. Pozzolanic
reaction occurs when lime and silica react in
the presence of water to form hydrated
calcium silicates. The degree of solids
modification is dependent on the final use for
the solids (e.g., road-base, concrete
aggregate, or structural fill). By-product
solids can be used as a road-base, concrete
aggregate, or structural fill. These
applications utilize improved properties of
FGD by-product mixed with fly ash:
increased unconfmed strength and decreased
permeability. These improved properties are
the result of pozzolanic reaction. Sometimes,
when commercial quality gypsum is made,
pelletization is employed. The selection of
the ultimate disposal method is highly site-
specific and depends on, among other factors,
land availability, hydrogeology, and
topography. In general, three options exist
for the ultimate disposal of waste FGD solids:
landfills, ponds, and gypsum stacks.
In addition to technical issues, several market
issues are involved in the decision of
wallboard manufacturers to use FGD
gypsum. These market issues are presented
below. Normally, the use of the quantity of
FGD gypsum produced by a representative
LSFO (hundreds of thousands of tons per
year) would be possible only if a dedicated
wallboard plant was built for this feed source,
or was shared by several existing wallboard
plants.8 The proximity of the wallboard plant
to the FGD by-product plant is important
because the transportation cost of FGD
gypsum to the wallboard plant can be a
significant percentage of its market value.
Since most existing wallboard plants in the
United States were designed to use mined
rock gypsum as feed material, the solids
handling equipment at these plants can use
only a limited quantity of FGD gypsum,
which has different handling properties.
Another potential obstacle to the
marketability of FGD gypsum is the fact that
the operating schedule of a power plant and
that of a wallboard plant often do not
coincide. Wallboard plants generally have
storage capacity to buffer the flow of gypsum
-------
in and out of the plant. Unlike power plants,
wallboard plants do not operate 24 hours per
day and 7 days a week. Similarly, power
plants do not operate year round, whereas
wallboard plants do.7
Limestone Inhibited Oxidation
A variation of the wet limestone process is
the limestone inhibited oxidation process
(LSIO). This process has been designed to
control oxidation in the absorber. The LSIO
is particularly well suited for applications
with high sulfur coals. Because of LSIO
chemistry, the difficulty in inhibiting the
oxidation generally increases with the
decreasing amount of sulfur content in coal.9
Several factors influence the performance of
LSIO. Flue gas composition, most notably
oxygen concentration, affects the extent of
sulfite oxidation to sulfate. Other flue gas
factors affecting LSIO are: SO2
concentration, fly ash content in the inlet gas
to the scrubber, and flue gas temperature and
humidity. Changing mass transfer
characteristics of the system (the ratio of SO2
/O2 absorbed) can alter the extent of natural
oxidation and, therefore, determine how
difficult it will be to inhibit the oxidation.
The change in mass transfer characteristics of
the system can result from adjusting the L/G.
Chemical characteristics of the system, such
as pH and liquid-phase composition, can also
alter the difficulty of oxidation inhibition.
In the LSIO, emulsified sodium thiosulfate
(Na2S2O3) is added to the limestone slurry
feed to prevent the oxidation to gypsum in the
absorber's internals by lowering the slurry
oxidation ratio to below 15 percent.10
Typically, a design oxidation ratio of between
4 and 10 percent is used in LSIO. The
amount of additive necessary to inhibit
oxidation depends on the chemistry and
operating conditions of a given absorber and
is, therefore, site specific.
Because of economic considerations, sulfur is
often added to the limestone slurry in lieu of
thiosulfate. Sulfur is added directly to the
limestone reagent tank. However, conversion
to thiosulfate occurs in the reaction tank when
sulfur contacts sulfite. The overall
conversion of sulfur to thiosulfate is between
50 and 75 percent. The amount of thiosulfate
(or sulfur) required to achieve inhibited
oxidation is a function of system chemistry
and operating conditions.
An additional benefit of using LSIO may be
an increased limestone solubility, which
enhances sorbent utilization. The waste
product, calcium sulfite, is landfilled. The
dewatering characteristics of the waste are
improved for LSIO compared to the waste
from natural oxidation operation of a wet
FGD absorber. This is because the calcium
sulfite product from the LSIO tends to form
larger crystals, similar to gypsum solids.
Lime and Magnesium-Lime
The lime process uses hydrated calcitic lime
slurry in a countercurrent spray tower. This
slurry is more reactive than limestone slurry,
but is more expensive. The magnesium-
enhanced lime process (MEL) is a variation
of the lime process in that it uses a special
type of lime, magnesium-enhanced lime
(typically 5-8 percent magnesium oxide) or
dolomitic lime (typically 20 percent
magnesium oxide).11 The operational pH
value for lime processes is normally in the
range 6.0 to 7.0 because of their increased
alkalinity and solubility, compared to
limestone processes. The lime process may
be designed to utilize the alkalinity of fly ash
in addition to the alkalinity of a sorbent.
10
-------
Lime used in the MEL contains magnesium
in addition to its calcitic component. Because
of the greater solubility of magnesium salts
compared to calcitic sorbents, the scrubbing
liquor is significantly more alkaline.
Therefore, MEL is able to achieve high SO2
removal efficiencies in significantly smaller
absorber towers than the limestone scrubbers.
Additionally, MEL allows for a significant
decrease of L/G, compared to LSFO for a
given target SO2 removal.12
Because waste solids from MEL have poorer
dewatering characteristics than solids from
calcitic limestone slurry processes, the best
dewatering operation of MEL occurs when
low solids concentration is maintained along
with moderate-to-low sulfite oxidation
levels.13 Forced oxidation, external to the
absorber, can be used in MEL to improve the
quality of their solids. This results in the
production of commercial quality gypsum.7
Commercial grade gypsum produced from
MEL is, in fact, brighter than gypsum
produced by a conventional LSFO. Brighter
gypsum, potentially, has a higher commercial
value.
14
Seawater Process
The seawater process utilizes the natural
alkalinity of seawater to neutralize SC>2. The
chemistry of the process is similar to the
LSFO chemistry except that the limestone
comes completely dissolved with the
seawater and that the chemistry does not
involve any dissolution or precipitation of
solids. Seawater is available in large amounts
at the power plant as cooling medium in the
condensers. It is used as a sorbent
downstream of the condensers for the purpose
of FGD. Seawater is alkaline by nature, and
has a large neutralizing capacity with respect
to SO2.
The absorption of SO2 takes place in an
absorber, where seawater and flue gas are
brought into close contact in a counter-
current flow. The scrubber effluent flows to
the treatment plant where it is air-sparged to
oxidize absorbed SO2 into sulfate before
discharge.15 The sulfate is completely
dissolved in seawater, so as a result there is
no waste product to dispose of. Sulfate is a
natural ingredient in seawater, and typically
there is only a slight increase of sulfate in the
discharge. This increase is within variations
naturally occurring in seawater. The
difference from the background level
normally is not detectable within even a short
distance from the point of discharge.
Since the utilization of seawater for SO2
scrubbing introduces a discharge to the ocean,
it is necessary to make an assessment based
on local conditions. Typically, the
assessment includes: effluent dilution and
dispersion calculations, description of
effluent, comparison of effluent data with
local quality criteria, description of local
marine environment, and evaluation of
possible effects from the discharge. High
chloride concentrations, characteristic of
systems using seawater, result in a
requirement for construction materials with
increased corrosion resistance.16
Dry FGD Technologies
In these technologies, SO2-containing flue
gas contacts alkaline (most often lime)
sorbent. As a result, dry waste is produced
with handling properties similar to fly ash.
The sorbent can be delivered to flue gas in an
aqueous slurry form [lime spray drying
process (LSD)] or as a dry powder [duct
sorbent injection process (DSI), furnace
11
-------
sorbent injection process (FSI), and
circulating fluidized bed process (CFB)].
The LSD and the CFB require dedicated
absorber vessels for sorbent to react with
SC>2, while in DSI and FSI new hardware
requirements are limited to sorbent delivery
equipment. In dry processes, sorbent
recirculation may be used to increase its
utilization. All dry FGD processes discussed
here are once-through (i.e., non-regenerable)
and, in general, limited to SC>2 removals
below those attainable with wet once-through
FGD.
Lime Spray Drying
LSD for the control of SC>2 emissions is used
for sources that burn low- to medium-sulfur
coal, with occasional applications for coals
with higher sulfur content. Some issues that
limit the use of spray dryers with high-sulfur
coals include the potential impact of chloride
contained in the coal on the spray dryer
performance, and the ability of the existing
paniculate control device to handle the
increased loading and achieve the required
efficiency.
The LSD is shown schematically in Figure 2-
3. Hot flue gas mixes in a spray dryer vessel
with a mist of finely atomized fresh lime
slurry. Fresh lime slurry is prepared in a
slaker (most often a ball mill) at a nominal
concentration of solids. Rotary atomizers or
two-fluid nozzles are used to finely disperse
lime slurry into flue gas. Typically, spray
dryers are operated at lime stoichiometry of
0.9 for low sulfur coals and 1.3 to 1.5 for high
sulfur coals. Simultaneous heat and mass
transfer between alkali in a finely dispersed
lime slurry and SCh from the gas phase result
in a series of reactions and a drying of
process waste. The amount of water fed into
the spray dryer is carefully controlled to
avoid complete saturation of the flue gas.
While a close approach to adiabatic saturation
(from 10 to 15 °C for coal-derived flue gas) is
required to achieve high SC>2 removal,
complete saturation impairs operation of a
spray dryer because of wet solids adhering to
vessel walls and within the particulate
collector. Primary reactions in the spray
dryer are as follows:
Ca(OH)2 + SO2 -» CaSO3 »/2H2O + l/2H2O (2-3)
Ca(OH)2 + SO3 + H2O -» CaSO4 • 2H2O
CaSO
CaSO
(2 - 4)
(2-5)
Some of the dry reaction product solids are
collected at the bottom of the spray dryer.
The remaining solids, suspended in the flue
gas, travel to the particulate control device
where the separation occurs. For a process
configuration where the particulate control
device is a baghouse, a significant additional
SC>2 removal may occur in the filter cake on
the surface of bags. Dry solids from the
particulate control device's hopper and from
the bottom of the spray dryer are disposed of.
The extent of alkali usage in a spray dryer is
limited by its available residence time for a
gas-solid reaction. Typical residence time in
a spray dryer is 8 to 12 seconds. In order to
increase sorbent utilization, part of the dry
solids from the bottom of the spray dryer and
the particulate collector's hopper are sent to
the recycle solids slurry tank. The
recirculated stream (shown with a broken line
in Figure 2-3) contains partially reacted alkali
from previous passes through the system.
The additional exposure of a sorbent to SC>2
afforded by the recycle promotes increased
sorbent utilization.
12
-------
Flue Gas
Quicklime
Water
Flue Gas
Out
Ball Mill
Slaker
W
Recycle Loop
Recycle Solids Slurry Tank
Chimney
Particulate
Control
Device
Disposal
Figure 2-3. Lime spray dryer FGD system.
-------
Mass transfer during a spray drying process
occurs in two discrete phases: moist and
dry.17 During the moist phase, SC>2 diffuses
from the bulk gas to the moisture layer on the
surface of lime particles and reacts with
dissolved lime. The reaction product
precipitates on the surface of the lime
particle. During the dry phase, 862 diffuses
through the products of the lime and 862
reaction and causes a gas-solid reaction with
the unreacted core of lime particle.
Studies indicated that a majority of 862
capture in the spray dryer occurs during the
moist phase. Any increase in the duration of
the moist phase would therefore increase the
amount of captured 862. Deliquescent salt
additives sometimes are added to the lime
slurry to be atomized in a spray dryer to
achieve this effect. A similar effect is
achieved when spray dryers are used on coals
with elevated chloride content.
Duct Sorbent Injection
DSI for 862 emission control is intended to
enable the control directly in the flue gas duct
between the air preheater and the particulate
control device. Since no dedicated absorber
vessel is required, the amount of hardware
needed to control SC>2 is minimized for DSI.
DSI utilizes the contacting of finely dispersed
sorbent with the flue gas. Sorbent used in
DSI is typically hydrated lime or,
occasionally, sodium bicarbonate.18 In the
DSI shown schematically in Figure 2-4,
dry hydrated lime sorbent is injected into
the flue gas downstream of the boiler's air
preheater. Water may be injected separately
from the sorbent either downstream or
upstream of the dry sorbent injection point to
humidify the flue gas. The relative position
of dry sorbent and water injection is
optimized to maximally promote the so-called
droplet scavenging or impacts between
sorbent particles and water droplets, both
suspended in the flue gas. Fly ash, reaction
products, and any unreacted sorbent are
collected in the particulate control device.
Additionally, recycling solids from the
particulate control device can boost the
utilization of alkaline material.19
A variation of DSI is duct spray drying
process, in which slurry is atomized and,
subsequently, evaporated in the duct.
Furnace Sorbent Injection
In the FSI, a dry sorbent is injected directly
into the furnace in the optimum temperature
region above the flame.20 FSI is shown
schematically in Figure 2-5. As a result of
the high temperature (approximately 1000
°C), sorbent particles (most often calcium
hydroxide, but sometimes calcium carbonate)
decompose and become porous solids with
high surface systems,21 according to the
reaction below:
Ca(OH)2
(2-6)
SC>2 in the flue gas reacts with the nascent
CaO as given below:
CaO + SO
CaSO4
(2-7)
Calcium sulfate, and any remaining unreacted
sorbent, leave the furnace with the flue gas.
In some systems, the flue gas is humidified
downstream of the air preheater or a
humidifier vessel is installed to improve
reagent utilization. Ex-situ spent sorbent
reactivation (wetting) is also used
occasionally as an integral part of the FSI.
Sorbent reactivated ex-situ is then injected
downstream of the air preheater. Such a
configuration should probably be considered
as a furnace/duct injection hybrid.
14
-------
Water
Boiler
Hydrated Lime
Silo
Figure 2-4. Schematic of DSI.
Flue Gas
Out
Chimney
Particulate
Control
Device
Disposal
-------
Sorbent
Injection
Humidifier
Reactivation
Reactor
Flue Gas
Out
Chimney
Particulate
Control
Device
Boiler
Disposal
Figure 2-5. Schematic of FSI.
-------
Circulating Fluidized Bed
In CFB, dry sorbent [most often Ca(OH)2] is
contacted with a humidified flue gas in a
circulating fluidized bed. CFB is shown
schematically in Figure 2-6. The fluidized
bed is formed as a result of flue gas flowing
upward through a bed of sorbent solids. The
CFB provides a long contact time between
the sorbent and flue gas because sorbent
passes through the bed several times. The
flue gas laden with reaction products then
flows to a particulate control device. Some of
the particulate control device's catch is
recirculated into the bed to increase the
utilization of sorbent, while the remaining
fraction is sent to disposal.
The CFB is characterized by good SO2 mass
transfer conditions from the gas to the solid
phase. This is achieved as a result of intimate
mixing of the solids with the gas as well as a
high slip velocity between the two phases.
An additional benefit of the fluidized bed is
continuous abrasion of sorbent particles,
resulting in the exposure of fresh, unreacted
alkali.
22
The CFB is not widely used in the United
States, and the bulk of its operating
experience comes from Germany for units
ranging from 50 to 250 MWe.23 This process
uses hydrated lime rather than the less
expensive and less reactive limestone
commonly used in wet FGD technology
processes. Additionally, due to a higher
particulate matter concentration downstream
of the fluidized bed, a larger ESP (or an
additional precollector) may be needed to
maintain the required particulate emission
levels compared with a non-circulating
sorbent.
Regenerable FGD Technologies
Regenerable FGD technologies discussed in
this section include four wet regenerable
processes (sodium sulfite, magnesium oxide,
sodium carbonate, and amine) and one dry
regenerable process (activated carbon).
These processes are characterized by their
product, a concentrated stream of 862. As
will be discussed in the following section,
regenerable FGD technology finds only
marginal application in the United States and
throughout the world. These processes have
a comparatively high O&M cost relative to
other FGD processes, and the return from sale
of the product does not offset a significant
portion of the increased process cost. Product
marketability may be a major problem.24 As
a result, some of the existing regenerable
FGD-technology-equipped units have been
converted to advanced limestone wet FGD.25
Wet Regenerable FGD
Sodium Sulfite
The sodium sulfite, or Wellman-Lord
process, absorbs SC>2 in a wet scrubber where
pretreated flue gas is contacted with sodium
sulfite solution. The product of the reaction
is sodium bisulfite liquor heavily loaded with
SC>2. The liquor is subsequently regenerated
in evaporators that crystallize sodium sulfite.
Concentrated 862 is suitable for sulfuric acid
production.
Magnesium Oxide
In the magnesium oxide process, 862 is
removed in a wet scrubber. In this process,
hydrogen chloride and hydrogen fluoride are
removed in a prescrubber. The magnesium
sulfite/sulfate product results from 862
absorption in a scrubber. The absorbed
product is dried and calcined in a kiln to
regenerate magnesium oxide. SC>2 captured
during calcination is suitable for sulfuric acid
production.
17
-------
Precollector
N/
Boiler
Flue Gas
Out
Chimney
Particulate
Control
Device
Disposal
Circulating
Fluidized Bed
Reactor
Fresh Sorbent
Figure 2-6. Schematic of CFB.
-------
Sodium Carbonate
In this process, SC>2 is contacted with a spray
of sodium carbonate solution. Products of the
reaction are sodium sulfite and sodium
sulfate, which are reduced to sodium sulfide.
Following the reaction of sodium sulfide with
carbon dioxide and water, sodium carbonate
is regenerated and hydrogen sulfide is
converted to sulfur.26
Amine
The amine process involves absorption of
862 with an aqueous amine absorbent. The
amine is regenerated thermally to release a
concentrated water-saturated SO2 stream. SO2
may then be treated by conventional
technologies to produce sulfuric acid.
Dry Regenerable FGD
Activated Carbon
The activated carbon process adsorbs 862 on
a moving bed of granular activated carbon.
Activated carbon is thermally regenerated to
produce a concentrated 862 stream. 862 may
then be treated by conventional technologies
to produce sulfuric acid.
19
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CHAPTER 3
TECHNOLOGY APPLICATIONS
Introduction
As discussed before, FGD technology
applications were reviewed based on the
information in CoalPowerS Database,
available from the International Energy
Agency's Coal Research Centre in London,
England and released in November 1998.27.
This database has not been modified or
otherwise amended. Findings of this review
are described below.
Historical Applications
Applications of FGD technologies over the
last three decades are shown in Figures 3-1
and 3-2 for the United States and the world,
respectively. In the United States, wet FGD
technology has dominated throughout the '70s
and early '80s with over 90 percent of the
overall installed FGD capacity. This same
period also saw a considerably high rate of
FGD installation: approximately 25,000 MWe
from 1976 through 1980. The mid-to-late
'80s saw a lower rate of FGD capacity
increase, compared to that of the '70s. It was
in the '80s that the first dry and regenerable
systems were installed. The early '90s saw a
slow increase of installed FGD capacity, in
wet and dry FGD technologies. A significant
increase of the FGD capacity occurred from
1994 through 1998. During this period, as
much as a 20,000 MWe increase was
accomplished, almost all of it in wet FGD.
No significant increase in regenerable FGD
capacity has taken place since the early '80s.
A somewhat different pattern for the rate of
application of FGD technology could be
observed throughout the world, as shown in
Figure 3-2. With approximately 30,000 MWe
of installed FGD capacity in 1980, the
capacity has been increasing at an
approximate rate of 100,000 MWe per
decade. Similar to the trend in the United
States, no significant increase in regenerable
FGD capacity has taken place worldwide
since the early '80s. Also, the rest of the
world has seen a smaller percent of dry FGD-
controlled capacity than the United States.
Since the wet FGD technology has
historically dominated both U.S. and
worldwide applications, it is of interest to
analyze application data in terms of specific
wet FGD processes. An illustration of U.S.
applications is presented in Figure 3-3. The
initial installed FGD capacity in the early 70's
was dominated by limestone processes.
Shortly thereafter, lime processes (lime and
MEL) were applied. The sodium carbonate
process was first applied in late '70s, and this
application has not seen any significant
growth through 1998. The growth of FGD
during the mid-to-late '80s, as well as the
early '90s, was almost entirely due to the
increase of the wet limestone process
capacity. From 1994 through 1998, there was
a step increase in the installed FGD capacity
with most of this being attributed to wet
limestone processes and the dolomitic lime
process in the United States.
20
-------
120000
100000 - -
80000 - -
60000 - -
40000 - -
20000 - -
0
DREGENERABLE
DDRY
• WET
1970 1972 1974 1976 1978 1980 1982 1984 1986 1988 1990 1992 1994 1996 1998
YEAR INSTALLED
Figure 3-1. Historical application of FGD technology in the United States.
-------
250000
200000 - -
150000 --
looooo --
50000 - -
DREGENERABLE
DORY
•WET
1970 1972 1974 1976 1978 1980 1982 1984 1986 1988 1990 1992 1994 1996 1998
YEAR INSTALLED
Figure 3-2. Historical application of FGD technology throughout the world.
-------
80000 -
70000 -
60000 -
50000 -
D Mag LIME
•LIME
D SODIUM CARBONATE
•LIMESTONE
1970 1972 1974 1976 1978 1980 1982 1984 1986 1988 1990 1992 1994 1996 1998
YEAR INSTALLED
Figure 3-3. Wet FGD technology application in the United States.
-------
Historical applications of dry technologies in
the United States are shown in Figure 3-4.
As presented in this figure, the spray drying
process has historically dominated
applications in the United States throughout
the '80s and '90s. The late '80s and early '90s
saw a mild increase of the installed capacity
of duct sorbent injection. There were also a
few furnace sorbent injection commercial
applications during the early '90s and CFB
applications in the mid '90s. Clearly, the
spray drying process has been popular among
the dry FGD technology processes.
Finally, historical applications of regenerable
processes in the United States are shown in
Figure 3-5. Regenerable processes (e.g.,
sodium sulfite, magnesium oxide, sodium
carbonate, activated carbon, amine) have not
seen any increase in their installed capacity
following their initial application.
In summary, the majority of historical
applications of FGD technology in the United
States, as well as throughout the world, have
utilized wet limestone and spray drying
processes. Wet FGD technology, other than
the wet limestone process, either uses a more
expensive sorbent (lime) or is limited by the
local availability of the specific sorbent used
by the process (e.g. sodium carbonate
process). Dry FGD technology, other than
LSD, either does not enjoy significant
commercial experience (e.g., CFB and FSI)
or offers only limited sorbent utilization (e.g.,
DSI).
The LSD has enjoyed a relatively steady
increase in installed capacity in the United
States since its initial application in the early
'80s. Wet limestone installed capacity
increased sharply during the '80s, stagnated
during early '90s, then experienced a step
increase during the late '90s (due to the
impact of the Clean Air Act Amendments of
1990).
Current Application
Table 3-1 shows statistics describing the
installation of FGD systems at fossil-fuel -
fired electric power plants through 1998.
FGD systems were installed to control 862
emissions from over 226,000 MWe
generating capacity, worldwide. Of FGD
systems installed on this capacity, 86.8
percent consist of wet FGD technology, 10.9
percent consist of dry FGD technology, and
the balance consist of regenerable FGD
technologies. Through 1998, almost 100,000
MWe of capacity in the United States had
FGD technology. Of these FGD systems
installed, 82.9 percent consist of wet FGD
technology, 14.2 percent consist of dry FGD
technology, and the balance consist of
regenerable FGD technologies. The percent
shares of the three FGD technology
categories installed are shown in Figure 3-6.
The pattern of installations in the U.S. and
abroad reflects that wet FGD technologies
predominate over other FGD technologies. It
is generally recognized that high SO2 removal
efficiency, coupled with cost effectiveness,
has been responsible for the overwhelming
popularity of wet FGD technologies,
particularly wet-limestone-based FGD
technologies. While the earlier wet FGD
systems produced only waste by-product
sludge, recent systems produce salable by-
product gypsum. This has likely increased
the attractiveness of wet FGD technologies.
Limited application of dry FGD technologies,
compared to wet FGD technologies, is likely
the result of their higher reagent cost and
limited choices for by-product disposal.
24
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16000
14000
12000
10000 -
8000 -
6000 —
4000 -
2000 --
D CIRCULATING FLUID BED
D FURNACE INJECTION
D SPRAY DRYING
• DUCT SORBENT INJECTION
1974 1976 1978 1980 1982 1984 1986 1988 1990 1992 1994 1996 1998
YEAR INSTALLED
Figure 3-4. Dry FGD technology application in the United States.
-------
ON
3000
DMgO
• SODIUM SULFITE
DOTHER
1980 1982 1984 1986 1988 1990 1992 1994 1996 1998
YEAR
Figure 3-5. Regenerable FGD technology application in the United States.
-------
K»
^1
U.S.
2.9%
82.9%
8.3%
Abroad
1.9%
89.8%
Worldwide
10.9%
2.3%
86.8%
fjDry
| [Regenerate
Figure 3-6. Percent shares (capacity) of the three FGD technologies installed.
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Table 3-1. Coal-fired Electrical Generation Capacity (MWe) Equipped with FGD Technology (1998)
Technology United States Abroad World Total
Wet
Dry
Regenerable
Total FGD
82,092
14,081
2,798
98,971
114,800
10,654
2,394
127,848
196,892
24,735
5,192
226,819
Table 3-2 shows capacities of various wet
FGD technology systems at power plants in
the United States and abroad. Of the United
States wet FGD technology installations, 68.9
percent use limestone processes. Abroad,
limestone processes are used on as much as
93.2 percent of the total wet FGD technology
installations. This trend is shown in Figure 3-
7, which shows the division of wet FGD
technology applications into limestone and
non-limestone ones. The main difference in
the pattern of wet FGD technology use in the
United States and abroad is the extent of the
application of dolomitic lime and sodium
carbonate processes. The attractiveness of
these processes depends on the local
availability of the special sorbents they
require. Limited availability of these special
sorbents abroad has likely limited the
application of the two processes. In the U.S.,
dolomitic lime and sodium carbonate
processes have been applied on some units
due to reagent availability at particular sites.
Table 3-3 shows statistics describing the
pattern of use of dry FGD technologies. Of
the worldwide capacity equipped with dry
FGD technology, 73.7 percent use the spray
drying process. This compares with 80.4
percent equipped with the spray drying
process in the U.S. Almost all of the
remaining installations of dry FGD
technology use sorbent injection, which
includes furnace (with and without a
downstream humidifier) and duct (calcium
compound as well as sodium compound)
injection. The dominance of the spray drying
process within the dry FGD technology
category is because this process is more
economical for low-to-moderate-sulfur coal
applications than wet FGD technology.
These processes have been used
commercially in the U.S. since the early '80s
and abroad since the mid '80s. Other dry
technology processes are considered to be
niche applications for retrofit systems, where
only limited SO2 removal is required.
Further understanding of recent FGD
technology selections made by the U.S.
electricity generating industry can be gained
by examining the recent FGD technology
installations in the U.S. Between 1991 and
1995, 19,154 MW of U.S. electric generating
capacity were retrofitted with FGD
technologies. Of this capacity 75, 17.5, and
7.5 percent were equipped with LSFO, MEL,
and LSD, respectively.
28
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Table 3-2. Total Capacity (MWe) Equipped with Wet FGD Technology (1998)
Process United States Abroad
World Total
Limestone 56,560 106,939
Lime 14,237 4,338
MEL 8,464 50
Sodium Carbonate 2,756
Seawater 75 1,050
Regenerable (other) - 2,423
Total Wet FGD 82,092 114,800
Table 3-3. Total Capacity (MWe) Equipped with Dry FGD Technology (1998)
Process United States Abroad
Spray Drying 11,315 6,904
Dry Sorbent Injection 2,400 1,125
CFB 80 517
FSI 286 2,108
Total Dry FGD 14,081 10,654
163,499
18,575
8,514
2,756
1,125
2,423
196,892
World Total
18,219
3,525
597
2,394
24,735
29
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U.S.
82,092 MWe
31.1%
Abroad
114,800 MWe
6.8%
68.9%
93.2%
Limestone
Non-Limestone
World
196,892 MWe
17.0%
83.0%
Figure 3-7. Comparison of limestone and non-limestone wet FGD applications.
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Table 3-4 shows additional statistics
describing the worldwide installation of FGD
systems on electric power plants. Through
1998, 668 FGD systems have been installed.
Of the installed FGD systems, 522 were wet
FGD technology, 124 were dry FGD
technology, and the balance consisted of
regenerable FGD technologies. Through
1998, 236 FGD technology systems were
installed in the U.S. Of the installed FGD
systems, 174 were wet FGD technologies, 54
were dry FGD technologies, and the balance
consisted of regenerable FGD technologies.
Combining the data from Table 3-4 with
those from Table 3-1 allows calculation of
representative sizes of FGD systems for each
of the technologies considered. These
representative sizes are shown in Table 3-5.
These average sizes were arrived at by
dividing the MWe shown in Table 3-1 by the
pertinent number of FGD systems shown in
Table 3-4.
As seen in Table 3-5, the installations of wet
FGD technology in the U.S., as well as those
abroad, appear to be larger than installations
of dry or regenerable categories of FGD
technologies. Additionally, the average FGD
system size in the United States is
considerably larger than abroad.
Table 3-4. Number of Installed FGD Technology Systems (1998)
Technology
Wet
Dry
Regenerable
Total FGD
United States
174
54
8
236
Abroad
348
70
14
432
World Total
522
124
22
668
Table 3-5. Average Size (MWe) of FGD Technology Systems (1998)
Technology
United States
Abroad
World Total
Wet
Dry
Regenerable
472
261
350
330
152
171
377
199
236
31
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CHAPTER 4
PERFORMANCE
Introduction
As discussed in Chapter 3, LSFO, MEL, and
LSD have been the dominant processes in
terms of the electric generating capacity
equipped with FGD over the last 30 years.
Therefore, the remainder of this report will
focus on issues related to these processes.
Removal Efficiency
An estimate of SO2 removal performance of
FGD processes can be obtained by examining
the design 862 removal efficiencies of these
processes reported in the CoalPowerS
Database. Table 4-1 shows design SO2
removal efficiencies for wet limestone and
LSD processes. These data reflect that wet
limestone systems have been designed for
high levels of SC>2 removal, up to 98 percent.
However, most wet limestone systems appear
to be designed for 90 percent 862 removal.
All LSFO systems installed after 1990 have
design SC>2 removal greater than 90 percent.
The units with low design efficiencies are
generally associated with plants burning low
sulfur fuels.28 Also, the units with the design
efficiency at the low end of the range given in
Table 4-1 are reported by the CoalPowerS
Database to have been installed in the 70s. It
is likely that the low design efficiencies are a
result of unit specific requirements for
permitting purposes, rather than technology
limitations. It is also likely that new
regulatory requirements were a catalyst for
technology improvements by creating a
market for more stringent SC>2 control.
Even though the median design efficiency for
all units with wet limestone processes in
CoalPowerS Database is 90 percent, it should
be emphasized that advanced, state-of-the-art
wet scrubbers are capable of routinely
achieving SC>2 removal efficiencies of over
95 percent. The high velocity LSFO process,
with state-of-the-art design options, is
reportedly capable of removing more than
99.6 percent of SO2 under test conditions.29
As seen in Table 4-1, the range and median of
SO2 reduction efficiency at LSD installations
are 70-96 and 90 percent, respectively. Spray
dryers often achieve greater than 90 percent
SO2 removal on coals with up to 2 percent
sulfur.30'31 CoalPowerS data also indicate that
all spray dryers installed during the period
from 1991 to 1995 have a design SO2
removal efficiency of between 90 and 95
percent.
The performance of wet limestone and LSD
processes has improved significantly over the
period of their application. To investigate
this improvement, the median design SO2
removal efficiency was determined for the
pertinent populations of wet limestone and
LSD installations for each of the three
decades: 1970-1979, 1980-1989, and 1990-
1999. The design efficiencies reported in the
CoalPowerS Database were used to determine
median design SO2 removal efficiency.
32
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Table 4-1. Design SO2 Removal Efficiencies
FGD Technology
Range of Design Efficiency,
percent
Median Design Efficiency,3
percent
Wet Limestone Processes
LSD Processes
52-98
70-96
90
90
"Derived based on CoalPower3 reported data. Application conditions for wet limestone and LSD processes may differ (e.g., coal sulfur percent).
Since the LSD did not become commercial
until the early '80s, no median efficiency
could be characterized for the '70s for this
process. For each of the last three decades,
median design SCh removal efficiencies, as
well as ranges of reported design SO2
removal efficiencies, for the wet limestone
and LSD are shown in Figure 4-1. A steady
increase of the design SO2 removal efficiency
can be noted for wet limestone and spray
drying processes. This improvement may be
due, in part, to the increasing need to better
control SC>2 emissions. However, the trends
do reflect that the SO2 removal efficiency for
the processes considered has improved with
time.
Energy Requirements
As described previously, once-through wet
FGD technology (and specifically, LSFO) has
enjoyed the largest extent of application
among all FGD technologies. Therefore, it
would be reasonable to expect any efforts
undertaken to improve energy efficiency of
FGD to be initiated on once-through wet
FGD systems. A review of the existing
literature reveals numerous efforts aimed at
increasing energy efficiency of wet FGD
systems. Both, design and operational issues
were considered in order to improve the
energy efficiency.
Modern LSFO absorbers operate at high flue
gas velocities in order to achieve improved
mass transfer and decrease absorber capital
cost at the same time. Flue gas velocity as
high as 20 ft/s was achieved under test
conditions. In an effort to improve the energy
efficiency, a new inlet design has been
implemented that incorporates the inlet
duct/absorber transition into the flared section
of the absorber. It is claimed that this new
design allows for a 33 percent pressure drop
reduction for absorbers operated at as much
as 20 ft/s gas velocity.
32
In a recent survey of LSFO O&M cost,33
pumping of sorbent slurry was consistently
ranked as the most energy intensive
component in the operation of wet FGD
systems. Pumping sorbent slurry raises the
slurry from tank to spray header level and
provides pressure necessary for fine
atomization. A decrease in the efficiency of
droplet/flue gas mixing must be compensated
for by increasing L/G in order to maintain the
target efficiency for SO2 removal. Therefore,
it is important to utilize a spray that has been
atomized within the spray tower for
maximum contact with the flue gas. In-depth
computational fluid dynamics studies,
coupled with field tests, have revealed a
radial gradient of SO2 concentration in a
33
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100
90
80
70
60 -
50
0
Wet Limestone
Spray Drying
Median
1970s
1980s
1990s
Figure 4-1. Design SO2 removal efficiencies for wet limestone and spray drying processes.
-------
wet limestone absorber.34 To remedy this
undesirable occurrence, guide vanes along the
perimeter of the tower could be used. When
installed on a 250 MWe absorber, the guide
vanes allowed for a 30-percent L/G reduction.
This reduction in L/G cut energy
consumption by as much as 20 percent.
Another energy intensive system in the
operation of LSFO system is limestone
pulverization. The quality and fineness of
grinding are critical operational parameters
that affect mass transfer properties in an
absorber. Horizontal ball milling is a
preferred method to pulverize limestone for
wet FGD. It is well suited to FGD service
because it offers a large reduction capability,
resistance to abrasion, and relatively low
operation, control, and maintenance
requirements.35 Depending on the mode of
grinding, a horizontal ball mill consumes 32
and 25 kWh/dry ton of limestone for the dry
and wet mode of operations, respectively.
Attrition grinding, a new method being
considered, has allowed for a reduction of
approximately 50 percent in energy
consumption, and uses only 15 kWh/dry ton
of limestone. An attrition grinder involves a
stationary vessel and internally stirred
grinding media (balls). Continuous attrition
grinders have been demonstrated that are
capable of grinding 6 mm limestone down to
95 percent minus 325 mesh.35
EPA's recently published cost estimation
algorithm, CUECost, estimates energy power
requirements for LSFO and LSD. CUECost
estimates energy consumption for LSFO
without DBA addition at 2 percent of the net
generating capacity of the unit prior to adding
pollution controls. With DBA addition, the
LSFO power consumption estimate is
reduced to 1.65 percent of the net generating
capacity. The LSD power consumption is
estimated at 0.7 percent of the net generating
capacity.
Applicability
There are some technical constraints to using
the spray drying process on applications with
high sulfur coal. In the U.S., this process has
typically been used in applications on units
burning low-to-medium-sulfur coal.36 There
has been a great deal of discussion regarding
the use of this process on units with high
sulfur coal requiring removal efficiencies of
over 80 percent. For each spray dryer, there
exists a maximum solids concentration
(sorbent slurry concentration) above which
the slurry cannot be easily atomized. High
sulfur coal applications may require sorbent
slurry concentrations in excess of the
maximum, since the amount of water that can
be evaporated is limited by the desired
approach to adiabatic saturation and
temperature of the flue gas leaving the
absorber.
Another technical constraint may be the
unit's physical size, which is a function of the
amount of flue gas to be treated. Typically,
spray drying has been applied to generating
units smaller than 300 MWe.36 However,
spray dryers have also been installed on
larger units using multiple absorbers.
Successful operation of a spray dryer is
dependent on a uniform mixing of finely
atomized sorbent slurry with flue gas. In
large spray dryer vessels, the limited
penetration of the atomized sorbent slurry
may compromise control efficiency.
-------
Once-through Wet FGD Technology
At present, several technical options exist for
upgrading the performance of existing
installations using wet limestone processes.
CHAPTER 5
ADVANCES
Introduction
Over the last 30 years, significant advances
have been made in wet limestone FGD
processes. As discussed before, once-through
dry FGD is a newer technology (applications
began in early '80s) and only a few
applications were seen in the United States
during the late '80s and during the '90s.
Since once-through wet FGD has been
involved with the bulk of FGD technology
applications during this period, no significant
advances in once-through dry FGD have been
reported. Therefore, only recent advances in
wet FGD will be discussed in this report.
Some of these advances have been aimed at
improving the performance and cost-
effectiveness of established processes, while
others have focused on developing new
processes. The initial part of this chapter
discusses once-through wet FGD technology
advances. It discusses both advances that can
be used to increase the performance of
existing once-through wet FGD systems and
advances that can be used in the construction
of new once-through wet FGD systems. The
chapter then concludes with discussion of a
new technology - ammonia scrubbing.
These options include:37
• increasing the sorbent amount used per
mole of SC>2;
• increasing the reactivity of the limestone
slurry with organic acid (e.g., dibasic
acid) addition;
• using more reactive sorbents;
• increasing L/G by increasing the recycle
slurry flow rate (requires more pumping
power);
• installing a perforated tray or other device
to increase mass transfer;
• reducing the amount of gas that is
bypassed (requires more fan power); and
• improving gas/liquid hydrodynamics (e.g.
guide vanes).
In general, selecting from the above options,
the existing installations may be upgraded to
achieve removal efficiencies of 95 percent or
more.
When considering the feasibility of upgrade
scenarios, interrelations between increased
SC>2 removal efficiency and many physical
and technical parameters require a thorough
evaluation. For example, the addition of
more sorbent may require the expansion of
the reagent preparation capacity and may
require better or increased sorbent preparation
(milling) capacity. Any increase in efficiency
will result in increased waste output, slurry
transport, dewatering, and waste disposal
capacity.38
The economics of FGD processes, affected by
technical advances and regulatory
36
-------
requirements, are driving numerous
conversions of existing older wet FGD
systems to more advanced ones. These
conversions are aimed at achieving improved
SC>2 removal efficiencies and/or waste
minimization. Limestone wet FGD systems
can be converted to MEL systems to increase
SC>2 removal efficiency. For example, an
inhibited oxidation limestone wet scrubber
designed for 85 percent SC>2 removal at an
L/G of 70 (gal/1,000 ft3) and 10 ft/s velocity
has been converted to MEL lime.39
Following the conversion, SC>2 removal
efficiency increased to 96.7 percent at an L/G
of 23 (gal/1,000 ft3).
In another example of a vintage wet FGD
system upgrade, conversion of an inhibited
oxidation wet FGD process to a LSFO system
was initiated in 1997.40 The objective of this
conversion was to initiate production of
commercial-grade gypsum in place of
calcium sulfite waste, which used to be
fixated via pozzolanic reaction with lime and
fly ash prior to disposal in a landfill.
Several advanced design, process, and
sorbent options are now available for new wet
FGD scrubbers.41 These options are shown in
Table 5-1. If implemented, some of these
advanced design options are capable of
providing high SC>2 removal and/or
improving the operational efficiency of wet
scrubbers while at the same time, reducing
cost.
Table 5-1. Advanced Options for New Wet FGD Scrubbers
Option
Design
Approach
large capacity modules
increased flue gas velocity in scrubber
concurrent flow
improved mist eliminator
improved hydraulics
superior materials of construction
low-energy spray nozzles
Sorbent
organic acid buffering
ultrafine limestone grind
Process
wet stack
in-situ oxidation
ex-situ oxidation with MEL
wastewater evaporation system
gypsum stacking for final disposal
37
-------
Among design improvement options,
construction of large capacity modules (single
module per unit) results in significant capital
savings (up to 35 percent) compared to the
baseline multi-module design.42'43 A single
tower absorber serving 890 MWe and two
units has recently been reported as being
under construction.44 The FGD system went
into operation in January 2000 as the largest
absorber in the United States and one of the
largest in the world.
Increased flue gas velocity in the scrubber
allows for a reduced vessel size. The reduced
vessel size is possible because of increased
mass transfer coefficients resulting from
higher gas/liquid relative velocity, increased
turbulence, and increased percentage of
droplets suspended within the
scrubber.12'29'45'46 Utilization of a concurrent
flow pattern provides a benefit in the form of
a reduced pressure drop across the vessel.
47
A considerable amount of computational fluid
dynamics modeling effort has been invested
in design advances for mist eliminators.48
Modifications include shape (forward tilt into
the gas flow), spacing (additional drainage),
and orientation (horizontal better than vertical
for high velocity scrubbers). These
modifications benefit the user with improved
mist eliminator cleaning, reduced
liquid/particulate matter carryover, and
minimized droplet re-entrainment.
Design modifications also include improved
hydraulics intended to intensify gas/liquid
contact throughout the system. Intensified
gas/liquid contact results in improved gas
velocity profiles across the spray tower.49
Improvements include: optimized placement
and selection of nozzles as well as installation
of wall rings to eliminate sneakage close to
the wall.29'50 Hydraulic model tests have
revealed that an optimum positioning of the
flue gas inlet in the flared section of the
absorber can significantly reduce its overall
pressure drop.32 In order to provide better
mixing of air and slurry as well as to improve
air distribution in the reaction tank, a rotary
air sparger has been used for LSFO.51
Finally, the advanced wet absorber system
design includes new materials of
construction, such as alloys, clad carbon steel,
and fiberglass, to provide corrosion resistance
at an optimum cost.52 In addition,
electrochemical protection is being used to
minimize the corrosion in reaction tanks for
systems with high fluoride concentrations.
This type of corrosion protection has been
determined as the most cost effective for such
applications.
53
Among improved sorbent options, the use of
organic acid buffering allows a reduced
vessel size and/or increased efficiency
through increased sorbent utilization. Organic
acids, such as dibasic acid (DBA), can be
added to limestone slurry in a wet limestone
process to improve SC>2 removal, sorbent
utilization, and/or a particular system's
operation. The increased 862 removal
efficiency in the presence of DBA is a result
of its buffering action (limiting the pH drop)
at the liquid/gas interface.54
An ultrafme limestone grind improves
limestone dissolution in the reaction tank
(reaction tank size reduction) and even in a
spray zone.29'55 An additional option is to
implement the direct use of pulverized
limestone, eliminating the need for on-site
grinding.
Some process modifications are aimed at
increasing the energy efficiency of the
process and include operation with a wet
38
-------
stack (no gas reheat) and a wastewater
evaporation system. The latter is
accomplished by liquid purge injection into
the hot flue gas upstream of the electrostatic
precipitator (ESP). In this option, wastewater
from solids handling/dewatering operation is
evaporated in the flue gas. Other process
options include in-situ forced oxidation,
which results in waste with better dewatering
characteristics for disposal.56
Recent process advances in MEL FGD
technology on full commercial-scale
incorporate ex-situ oxidation to produce
gypsum with excellent purity and bright
white color. By-product Mg(OH)2 can be
produced optionally for in-plant use or sale.
This Mg(OH)2 can be used for boiler
injection for SOs control, to minimize air
preheater fouling, and/or PM2.5-related stack
emission.
MEL can offer some advantages over LSFO.
It can operate with high 862 removal
efficiency (98 percent plus) in high sulfur
coal applications, low L/G ratio, smaller
scrubbers and recirculating pumps, and lower
energy requirement.
Ammonia Scrubbing
Over the last few years, a promising wet FGD
process has been under development. This
process, wet ammonia FGD, has the potential
to improve waste management in conjunction
with providing SO2 removal efficiency in
excess of 95 percent.57 Operators of
conventional wet limestone FGD processes
may be confronted with saturated markets for
commercial-grade gypsum of FGD origin. At
present, the wet ammonia FGD process offers
the advantage of an attractive ammonium
sulfate [(NFL^SO/t] by-product that can be
used as fertilizer.
This process also has the potential for
becoming a promising option for units
burning high sulfur coal, as it is also capable
of removing other acid gases (e.g., sulfur
trioxide [SOs] and hydrogen chloride [HC1])
in addition to SO2. While HC1 emissions can
be reduced concurrently with SO2 emissions
using currently commercial FGD technology,
the removal of SOs and control of sulfuric
acid (H2SO4) aerosol is not as
straightforward. Depending on the type of
FGD technology, a considerable portion of
H2SO4 aerosol may exit the stack as a
respirable fine particulate emission and may
cause a visible plume.58
Ammonia scrubbing of SO2 offers an
alternative for maximizing the value of the
by-product produced in a wet FGD system.59
With the ongoing deregulation in the electric
utility industry, the cost of generation for
large power generation units is continually
under scrutiny. For units that utilize wet,
limestone-based FGD, producing a salable
by-product, such as gypsum, is a means for
reducing the cost of operation. However, the
United States has an abundant supply of
natural gypsum and, as a result, the price for
FGD gypsum produced by LSFO system has
historically been very low.60 Apparent
problems related to the economics of FGD
gypsum can potentially be overcome by
ammonia scrubbing. The reaction of
ammonia, SO2, and oxygen in an absorber
installed in an ammonia scrubbing system
produces (ML;)2SO4 fertilizer. Recently,
prices of ammonia are reported to have
decreased considerably,60 making the use of
ammonia as a reagent much more
economically favorable compared to a few
years ago. The best opportunities to apply
ammonia scrubbing technology will likely be
found at power plants in a proximity to
39
-------
navigable water or good rail access, and a
location with high (NEL^SC^ prices.
The ammonia scrubbing process, as currently
envisioned,61 employs a counterflow spray
tower design that is similar in configuration
to the existing wet limestone-based FGD
systems. In some cases, prescrubber may be
used to humidify the flue gas and/or remove
HC1 prior to the main absorption stage. The
flue gas then enters the counterflow spray
tower where it is contacted by a solution of
(NH4)2SO4 liquor.
The ammonia is stored in a pressurized or
refrigerated vessel and pumped as a liquid to
a vaporizer. The vaporizer typically uses
steam to vaporize the ammonia prior to
introducing it into the oxidation air or directly
into the absorber reaction tank.
Ammonia is added with the oxidation air to
maintain the recycle liquor at the desired pH
to ensure that the required 862 removal is
achieved. The cleaned flue gas passes
through mist eliminators to remove any
entrained droplets. The absorber is operated
in a pH range selected to eliminate ammonia
slip and aerosol formation. Conversion of
ammonium sulfite [(NFL^SOs] and
ammonium bisulfite (NFLjHSOs) to
(NH4)2SO4 takes place in the reaction tank via
injection of compressed air. (NFL^SC^
solution (10 - 25 wt percent dissolved solids)
is bled from the absorber. Fresh makeup
water required by the process is added to the
absorber reaction tank to maintain the liquid
level in the tank.
61
Aerosol emissions are a concern with
ammonia scrubbing processes and must be
addressed carefully in the design. The
simultaneous presence of ammonia, SCVSO
and water vapor in flue gas can result in the
formation of ammonia/sulfur aerosols. The
aerosols are very small (0.1 to 0.3
micrometer) and, once formed, are emitted
from the absorber, causing a visible plume at
the stack discharge.
Although it is theoretically possible to
operate an ammonia scrubbing system in a
plume-free mode with very precise control of
absorber pH, temperature, solution
concentration, etc., wet electrostatic
precipitators (WESPs) are generally
employed downstream of the absorber to
eliminate concerns related to aerosol
emissions. Even with excellent process
control, aerosol emissions can occur as a
result of load changes, sulfur inlet changes,
non-ideal gas/liquid contact, and pH control
problems. An added benefit of the WESP is
the control of 863 mist and droplet emissions.
2SO4 fertilizer production can be
accomplished in two ways:
• Remote crystallization and drying
• In-situ crystallization, dewatering, and
granulation
In the case of remote crystallization/drying,
the absorber loop operates with a clear
solution of (NH4)2SO4 (approximately 30 to
35 percent). The solution is sent to an
adjacent by-product processing plant, which
consists of a crystallizer, centrifuge, and
dryer. Thermal energy (steam) or vapor
recompression/evaporation is used to
concentrate the solution to the point where
crystallization takes place.
40
-------
Advantages to clear solution operation
include:
• Discreet (NEL^SC^ crystals are formed in
a device specifically designed for that
purpose, so that the size can be carefully
controlled to meet the required product
specifications
• The monolithic crystals are not subject to
attrition or dusting during shipping or
handling
• The (NH4)2SO4 solution can be filtered
prior to crystallization, thus eliminating
any concern with solid contaminants (e.g.
fly ash) in the byproduct
• The entire absorber loop operates with
clear solutions and is not subject to the
plugging and erosion concerns associated
with slurry scrubbing
With in-situ crystallization processes, slurry
from the pre-absorber is passed to a
dewatering hydroclone, where the slurry
solids concentration is increased to about 35
weight percent. The purpose of the
hydroclone is two-fold: to dewater the slurry
from the prescrubber to optimize the
centrifuge feed slurry density; and to separate
the fine particles (primarily ash from the
boiler) from the product, and thus maintain
product purity. The slurry is next pumped to
a series of centrifuges where the slurry is
dewatered to 97 - 98 percent solids.
Centrifuges discharge the material
immediately into a rotary drum dryer where
heated air is passed over the crystals to
further dry the material to less than 1 percent
moisture.61
raw (NH4)2SO4 material is transferred to a
compaction system. In this system, fresh feed
2804 material is mixed with the raw
2804 in a pug mill mixer. Finally, the
material from the mixer is compacted into
hard flakes subsequently discharged into a
flake breaker. The flake breaker crushes the
large flakes into smaller pieces, later sized in
a series of sizing mills. The final acceptably
sized product is transported to the storage
area.
The chemistry of the production of
(NH4)2SO4 from boiler flue gas is very
similar to the chemistry of wet limestone
FGD. SC>2 from the flue gas is absorbed in the
spray tower by water according to the
equation:
SO2 + H2O H2SO3 (5-1)
The H2SO3 is then reacted in a reaction tank
with ammonia to form (NH4)2SO3 and
NH4HSO3:
H2SO3+2NH3 <-> (NH4)2SO3
H2SO3 + (NH4)2SO3 <-> 2NH4HSO3
(5-2)
(5-3)
(NH4)2SO3 and NH4HSO3 are also oxidized
in the absorber (forced oxidation) to form
(NH4)2SO4 and NH4HSO4:
(NH4)2S03 +/202 <-» (NH4)2S04
NH4HSO3
NH4HSO4
(5-4)
(5-5)
To maximize the by-product value, the
(NH4)2SO4 material must be converted to
larger granular crystals. To accomplish this,
41
-------
The NH4HSO4 is neutralized in the presence
of ammonia and water to form
NH4HSO4 + NH3 + H2O <->
(NH4)2SO4+H2O (5-6)
Because of the relatively high value of
(NH4)2SO4 fertilizer, the economics of
ammonia scrubbing improve as the sulfur
content of the fuel increases. Consequently,
ammonia scrubbing offers a potential to the
plant to use high sulfur fuel such as high
sulfur coal or petroleum coke. Thus, the need
for use of more expensive low sulfur coal
could be avoided. Petroleum coke has been
identified as a low cost, high sulfur fuel that
can be burned in many boilers. A recent
study61 concluded that, as more refineries use
crude with higher sulfur content, both the
quantity and sulfur content of the coke will
increase. Since many refineries are expected
to have difficulty disposing of the coke with
the high sulfur content, this fact should lead
to attractive prices for the material.
The attractiveness of the ammonia scrubbing
process appears to depend on the ability of
the plant to sell (NH4)2SO4 fertilizer. An
evaluation of the price of (NH4)2SO4 over a
period of 1 1 years has indicated a sustained
increase.61 This has been explained by its
value as a nutrient for selected crops and its
ability to replenish the sulfur deficiency in
soils.
A successful demonstration of 90-95 percent
SC>2 removal and aerosol-free operation has
recently been reported60 for a 130 MWe
system installed on boilers burning 2 to 3.5
percent sulfur coal.
42
-------
CHAPTER 6
FGD COST
General Approach
As discussed before, LSFO, LSD, and MEL
have been the processes of choice in recent
U.S. applications. Therefore, in this work,
state-of-the-art cost models were developed
for these processes. These state-of-the-art
models are collectively called State-of-the-art
Utility Scrubber Cost Model (SUSCM) and
are expected to provide budgetary cost
estimates for future applications. In the
ensuing paragraphs, descriptions and results
are provided for the state-of-the-art LSFO,
LSD, and MEL cost models developed in this
work.
The Air Pollution Prevention and Control
Division (APPCD) of EPA's National Risk
Management Research Laboratory (NRMRL)
has recently published the Coal Utility
Environmental Cost Workbook (CUECost).62
CUECost provides budgetary cost estimates
(+30 percent accuracy) for between 100 and
2000 MWe net LSFO and LSD applications
based on user-defined design and economic
criteria. CUECost algorithms provided the
starting point for the LSFO and LSD cost
models developed in this work.
For each of these models, first, a sensitivity
analysis was conducted to determine those
variables that have a minor impact on cost
(i.e., a deviation of less than 5 percent over
the selected baseline). Then, these variables
were fixed at typical values to arrive at a
simplified cost model. Next, the simplified
LSFO and LSD cost models were validated
with published data. Finally, these models
were further adjusted with cost-effective
design decisions to arrive at state-of-the-art
LSFO and LSD cost models.
For costing purposes, MEL can be considered
to be a combination of LSFO and LSD. In
the MEL, sorbent (magnesium-enhanced
slurry) is prepared in a similar manner to that
used in LSD, and this sorbent is contacted
with flue gas in an absorber similar to a
typical LSFO absorber. However, because
MEL sorbent is more reactive than LSFO
sorbent, less flue gas residence time is needed
in the MEL absorber. As such, a MEL
absorber is significantly smaller than a
corresponding LSFO absorber. Further MEL
waste handling equipment operates in a
fashion similar to that in LSFO, producing
gypsum by-product. Considering these
characteristics of the MEL, for costing
purposes this process can be considered to be
a combination of LSFO and LSD. Therefore,
the LSFO and LSD algorithms developed as
described above were used appropriately to
develop the MEL cost model. As for LSD
and LSFO, cost-effective design choices were
made to arrive at a state-of-the-art MEL cost
model.
Limestone Forced Oxidation
For the sensitivity analysis, the baseline
consisted of an LSFO application on a 500
MWe unit with a 10,500 Btu/kWh heat rate,
burning 3.4 percent sulfur (S) Jefferson, OH,
coal (heating value of 11,922 Btu/lb), and
presenting medium retrofit difficulty.
43
-------
The primary design elements fixed in this
baseline LSFO application were materials for
construction of the absorbers, addition of
DBA, a wet stack, and gypsum stacking
disposal. The choice of materials for
construction is known to have a major cost
impact. Selection of rubber-lined carbon
steel (RLCS) had the largest cost impact,
saving nominally 0.65 mills/kWh over alloy
construction. Other variables were fixed to
CUECost default values for the baseline
LSFO, including 95 percent removal of SO2.
This baseline LSFO retrofit requires one
absorber serving a maximum size of 700
MWe. Thus defined, baseline LSFO has an
annual operating cost of 10.31 mills/kWh.
Sensitivity analyses were performed using the
CUECost outputs resulting from single-
variable perturbations from the baseline. The
results of these analyses are summarized in
Table 6-1. Perturbations in the variables
were selected to span the range of realistic
values (e.g., unit size ranged between 100 and
2000 MWe). The high and low values of
variables were selected and the corresponding
costs were then determined for each single-
variable perturbation. Next, the prediction
differences were calculated between baseline
and high, as well as low, values for each
perturbed variable.
Based on the results of sensitivity analyses,
shown in Table 6-1, it was determined that
the majority of cost impacts (cost impacts
greater than + 5 percent) can be captured with
capacity, heat rate, coal sulfur content, coal
heating value, capacity factor, and disposal
mode.
The remaining variables were determined to
have a minor impact on cost and, therefore,
they were fixed at typical values. The list of
variables that have minor impacts on the cost,
as predicted by the sensitivity analyses, is
given in Table 6-2. Furthermore, the values
selected to fix these minor variables are also
shown in Table 6-2. For example, the air
heater outlet temperature that was shown to
have between 1.4 and -0.5 percent impact
when varied between 360 and 280 °F
respectively, was fixed at 300 °F, as shown
in Table 6-2. These fixed values are based on
the CUECost defaults.
Fixed operation and maintenance (O&M) cost
in the simplified LSFO cost model accounts
for the cost associated with operating labor,
maintenance labor and materials, and
administration and support labor. The
variable O&M cost is composed of reagent
cost, disposal cost, steam cost, and energy
cost. The assumptions used in calculating
these costs are based on the default values
provided in CUECost and the suggested
values in Electric Power Reserach Institute's
Technical Assessment Guide (EPRI TAG).
CUECost determines capital cost for FGD
system as Total Capital Requirement (TCR).
The cost estimation begins with the installed
equipment capital cost (BM). Following the
EPRI TAGs methodology, the installed BM
cost is then multiplied by appropriate factors
to incorporate costs of general facilities,
engineering fees, contingencies, and the
prime contractor's fee, resulting in an
estimate of Total Plant Cost (TPC). Financial
factors related to the time required to
construct the FGD equipment are applied to
TPC to estimate Total Plant Investment (TPI).
TCR is the sum of TPI, inventory cost, and
pre-production costs. Pre-production cost
incorporates one-twelfth of the projected
annual O&M expenses and 2 percent of the
TPI estimate.
44
-------
Table 6-1. Sensitivity Analysis of LSFO Annual Operating Cost (baseline cost of 10.31 mills/kWh)
Variable, units
Capacity,
MWe
Heat Rate,
Btu/kWh
Coal Sulfur
Content, %
Coal Heating
Value3, Btu/lb
Air Heater
Outlet, °F
SO2 Removal,
L/G
Slurry
Concentration,
% solids
Capacity
Factor, %
DBAC
Addition
Disposal Mode
Absorber
Material
No. of
Absorbers
Reheat
Baseline
500
10,500
3.43
11,922
300
95
125
15
65
no
stacking
alloy
1
yes
Variable's
High Value
2000
11,000
4.0
14,000
360
98
160
20
90
N/Tb
landfill
N/T
2
N/T
Variable's Cost for Cost for Low High Value Low Value
Low Value High Value Value of Difference, d Difference, e
of Variable, Variable, % %
mills/kWh mills/kWh
100
8,000
1.5
10,500
280
90
60
10
40
yes
wallboard
RLCS
N/T
no
6.57
12.25
10.60
9.56
10.45
10.36
10.36
10.31
8.04
N/T
12.51
N/T
10.41
N/T
22.62
9.13
9.71
10.92
10.26
10.22
10.22
10.31
15.41
10.20
10.23
9.66
N/T
9.94
-36.3
18.8
2.8
-7.3
1.4
0.5
0.5
0
-22.0
N/T
21.3
N/T
1.0
N/T
119.4
-11.4
-5.8
5.9
-0.5
-0.9
-0.9
0
49.5
-1.1
-0.1
-6.3
N/T
-3.6
"Coal Data: Form EIA-767, DOE 2S
bN/T=not tested
'Dibasic Acid
""Difference = (Cost for High Value of Variable - 10.31) /10.31 • 100%
'Difference = (Cost for Low Value of Variable - 10.31) /10.31 • 100%
45
-------
Table 6-2. Representative Values for LSFO Variables with Minor Cost Impacts
Variable
Coal Heating Value3
Limestone Composition
SO2 Control Efficiency
L/G
Ambient Pressure
Air Heater Outlet Temperature
Moisture in the Flue Gas
Max Fan Capacity
Chimney Inlet Gas Temperature
Units
Btu/lb
% CaCO3
%
gal/1,000 ft3
in. Hg
op
%
cfm
op
Value
11,900
95.3
95
125
29.4
300
14.0
1,600,000
127
Comments
Baseline 11,922
70 with DBA
Either 2, 4, or 8 fans
a Not a minor impact; value is set to 1 1,900 Btu/lb.
Capital Cost
LSFO systems consist of five major
equipment areas: reagent feed, SO2 removal,
flue gas handling, waste handling, and
support equipment. As described before,
capital cost algorithms for these areas in
CUECost were simplified to be functions of
capacity, heat rate, coal sulfur content, and
coal heating value only. Summation of these
adjusted algorithms provides the total capital
cost in the simplified LSFO cost model.
The above five areas are shown schematically
for LSFO in Figure 6-1. Accordingly, in cost
considerations the capital cost of each area is
represented as: Reagent Feed (BMp), SO2
Removal (BMR), Flue Gas Handling (BMG),
Waste Handling (BMW), and Support
Equipment (BME). The estimation methods
used for the five major equipment areas are
described below.
The BMp, BMR, and BMw cost estimates
were explicitly determined by the SO2 feed
rate to the FGD system. This feed rate was
determined by the coal sulfur content and
coal use rate, with no provision for sulfur
retention in the ash. SO2 flow rate to the
FGD system (FRso2) was estimated from the
amount of sulfur in the coal as well as the
coal burn rate at full load:
FRso, =
Wt%S»\000 ^
HHV
64
32
where Wt%S is coal sulfur content (wt%),
MWe is LSFO size (MWe), HR is plant heat
rate (Btu/kWh), and HHV is coal heating
value (Btu/lb).
Reagent Feed Area
The BMF cost (including receiving, storage,
and grinding) - a fourth order polynomial in
limestone addition rate was used based on
CUECost. The limestone addition rate was
determined based on the SO2 feed rate to the
absorber, reagent addition rate, SO2 removal
requirement, and limestone CaCOs content.
In CUECost (and in this simplified model),
all the sulfur in the coal was assumed to be
delivered to the FGD system as SO2.
46
-------
FLUE GAS HANDLING
WASTE/BY-PRODUCT
Figure 6-1. Schematics of LSFO system's equipment areas.
-------
CUECost adjusts the reagent feed ratio to
ensure the CaCOs present is sufficient to
remove all the chlorine in the coal as CaCb,
in addition to the specified 862 removal.
However, chlorine removal has been
eliminated from this model based on the
assumption that it has a negligible cost
impact. Specifically, the cost of limestone
and reagent addition was calculated as
follows:
- Cost of ball mill and hydrocyclones -
second order polynomial on limestone
addition rate.
Cost of DBA supply tank - power law
on DBA addition rate, which is, in
turn, proportional to the rate of 862
removal.
The BMp cost was estimated based on the
limestone feed rate. Limestone composition
(purity) has been fixed in this model at 95.3
percent CaCOs, which is the default
composition used in CUECost. The
limestone addition rate has been fixed in this
model at 1.05 times the reagent feed ratio.
Reagent feed rate (FRiJ was estimated as:
inn n CK
(6-2)
64 0.953
These parameters allow BMF to be estimated
as follows:
BMF = - 0.0034 <
^lOOOj
2.1128
UoooJ
I ( 7^7? i 1 7^7?
- 494.55 • - M + 68164.7 • - ^ + 7118470
^ UoooJ J 1000
+ CB&H + CDBA (
where CB&H is the cost of the ball mill and
hydroclones as given by:
- 3)
- +
ooo
ooo
1854902
+ 22412
(6-4)
and CDBA is the cost of the DBA tank as given
by:
CDBA= 364627.
FR
so,
0.95*20
2000
\0.283
8.34* (1 + 0.5)
60
(6-5)
CDBA was added only for LSFO systems with
the DBA addition.
SO? Removal Area
BMR cost (including absorbers, tanks, and
pumps) - a third order polynomial on 862
rate to the scrubber was used based on
CUECost. These cost components were
calculated as follows:
Cost of absorbers - power law on flue
gas flow rate to each absorber inlet
multiplied by the number of
absorbers. Different power laws were
used depending on absorber
construction materials. Maximum
absorber size was limited in CUECost
to treat 700 MWe; larger units
required multiple, equal size
absorbers.
Cost of spray pumps - power law
applied to the slurry flow rate per
absorber per pump multiplied by the
number of pumps. The slurry flow
rate (gpm) was calculated based on
the gas flow rate per absorber at the
exhaust temperature, but at 1 in. H2O
less than the inlet pressure (typical
absorber inlet pressure drop). L/G
48
-------
was fixed at 70 for LSFO with DBA
and, otherwise, was fixed at 125.
CUECost default is L/G of 125 for 95
percent 862 removal.
cost required estimation of flue gas flow
through the LSFO system. The absorber cost
was estimated based on inlet flue gas flow
rate and type of construction materials. The
spray pump cost was estimated based on flue
gas flow rates exhausting the absorber.
The flue gas flow was calculated in CUECost
using the coal analysis in addition to unit size,
heat rate, excess air, and air inleakage. This
approach was analogous to computing the F-
factor (Fd) for each fuel. As it was not
considered practical to calculate an Fd, for
each fuel, gas flow was estimated using the
methodology employed in 40CFR75
Appendix F. An Fd of 9,780 scf/106 Btu was
applied for all coals, as the differences in coal
rank (e.g., 9,860 scf/106Btu for lignite) were
expected to have negligible impact on the
estimated scrubber cost. Flue gas flow into
the absorber (ACFM) was calculated as
follows:
ACFM =
1000 9780 (460 + 295) 100
„ OT
a »HR»
106 60 528 (100-6)
0.04 0.209 (P-0.04)
P + P * (0.209-P)
(6-6)
where P is %C>2 in the stack (9 percent C>2 in
the stack was assumed).
The pressure at the absorber inlet was fixed at
10 in. H2O gauge, the CUECost default.
Ambient pressure was fixed at the CUECost
default of 29.4 in. Hg. Temperature of the
flue gas entering the absorber might have
been varying significantly for different units
but was expected to have minimal impact on
cost, based on the sensitivity analysis.
Absorber inlet temperature was fixed in the
model at 295 °F, resulting from 300 °F air
heater outlet temperature used as the default
in CUECost. The moisture fraction was
assumed to be 6.0 percent H^O at the
absorber inlet.
The cost of the spray pumps for the absorbers
was estimated based on the absorber outlet
flue gas flow rate and the number of pumps
(Np) required. The Np required was based on
the required slurry flow rate per absorber and
a maximum single pump capacity of 43,000
gpm (CUECost default). The required slurry
flow rate was determined by L/G, dependent
on whether the design incorporated DBA
additive. The gas flow rate was determined at
127 °F and at 9 in. H2O gauge (CUECost
default). Moisture content was estimated at
14 percent H2O (CUECost calculated).
CUECost estimated air addition at 2 moles
oxygen for each mole of sulfite to be oxidized
(CUECost default). For a typical SO2
concentration, this air addition is less than 1
percent by volume of the total flow and has
not been included.
The above assumptions allowed estimation of
BMR cost, depending on the absorber
construction material used and on the
presence of DBA addition in the system as
follows:
BMR = BAREMOD\JLER + ABSORBER
(6-7)
L is the absorber cost
equal
where ABSORBER is
to:
(ACFMT515
ABSORBER 1 = 173978. • Na (6-8)
i, 1000 J
or to:
49
-------
O.5638
ABSORBER 2 = 230064 • —
1000 J
for the RLCS or alloy material of
construction, respectively (Na is number of
absorbers).
The cost of pumps, PUMPS, was expressed
as:
(6-9)
PUMPS = 910.85 •
F,
GPM
•Nr
(6-10)
where GPM is slurry flow rate (gpm) and Np
is the number of pumps. The slurry flow rate
varied, depending on whether dibasic acid
additive was selected.
Auxiliary cost for the 862 Removal Area
(BARE MODULER) was calculated as
follows:
BAREMODULEp = 0.8701-
FR.
-so,
1000
-188.2
FR
-so,
1000
+ 34809 •
FR
-so,
1000
M +1905302
(6-11)
Flue Gas Handling Area
The flue gas handling system cost (ductwork
and ID fans) was based on CUECost, a
polynomial on flue gas flow rate entering the
absorbers, exiting absorbers, and number of
absorbers. If a design included reheat, a term
was added for the required temperature
increase. The cost of ID fans was estimated
using a power law based on the inlet gas flow
rate per fan multiplied by the number of fans
required.
The cost of the BMG was based on the
number of absorbers, flow entering the
absorbers, and flow exiting the absorbers.
Pressure of the gas exiting the absorbers was
fixed at 4 in. H2O gauge. The temperature of
the gas exiting the absorbers was fixed at 127
°F, the CUECost default wet bulb
temperature. Flue gas moisture content was
approximated at 14 percent H2O at the
absorber outlet and through the remainder of
the FGD system (CUECost default).
The cost of fans was estimated by a power
law based on the number of fans required and
the flue gas flow rate. Fans were assumed to
be installed in groups of 2, 4, or 8 with a
maximum individual fan capacity of
1,600,000 cfm. The number of fans was
based on conditional tests of the smallest
number option (2, 4, or 8) resulting in an
individual fan capacity of less than 1,600,000
cfm. Inlet pressure for sizing fans was fixed
at 12 in. H^O vacuum. Temperature at the
fan inlet was fixed at 295 °F (CUECost
method).
If the design incorporated reheat, BMo cost
was adjusted according to the design
temperature increase. The reheat temperature
rise was fixed at 25 °F (CUECost default).
By assuming the above design criteria,
cost was estimated as follows:
BMG = BARE MODULEG + ID FANS
(6-12)
50
-------
The auxiliary cost of the Flue Gas Handling
Area (BARE MODULEG) was calculated as:
BARE MODULE,, = -0.1195 •
ACFM}2
1000 J
( ACFM \
+ 777.76 • - + 238203 + 0.000012
i, 1000 )
1000 )
1000 )
.82
i, 1000 J
+ 1266.4.
+ 559693- 0.2009.
ACFM1
1000 *Nn
1000•Na
+ 420141 (6-13)
where ACFM1 is flue gas flow rate out of the
absorber.
The cost of fans (ID FANS) was calculated
as:
-.0.6842
ID FANS
= 91.24.U^ .N,
(6-14)
where Nf is the number of fans.
Waste/By-product Handling Area
The BMw cost (dewatering, disposal/storage,
and washing) - a second order polynomial on
SC>2 mass flow rate for gypsum stacking was
used based on CUECost. Moreover, a third
order polynomial on SC>2 mass flow rate was
used for landfill disposal or wallboard
gypsum production. The cost of thickener
was estimated as a linear function of the
waste solids removal rate. The waste amount
was estimated from a mass balance.
The BMw cost was fixed by the disposal
option chosen and by the amount of sludge to
be disposed of. The amount of sludge was
based on inlet 862 flow rate, 862 removal
efficiency (fixed at 95 percent), and CaCOs in
the limestone. All SC>2 removed was
assumed to be oxidized to form calcium
sulfate dihydrate (gypsum). The BMW cost
was estimated as follows:
BMW = BARE MODULE^ + THICKENER (6-15)
• For the Waste/By-product Handling
System with gypsum stacking (BARE
MODULEwi):
BMWI = -4.0567 •
FR.
-so2
1000
+ 1788
FR
•so.
1000
+ 80700
(6-16)
• For the Waste/By-product Handling
System with landfill (BARE
MODULEw2):
BMW2 = 0.325 •
f FR
+ 29091. - S°
FR,
1000
1000
+ 773243
-168.77 •
FR.
-so2
1000
(6-17)
• For the Waste/By-product Handling
System with wallboard gypsum
production (BARE MODULEws):
BMW3=BMW2»1.25 (6-18)
The cost of thickener (THICKENER) was
estimated as:
THICKENER = 9018.7 • FRvn • 0.95
•so,
172
64 •2000
-+114562
(6-19)
51
-------
Support Equipment Area
The BMs cost (electrical, water, and air) - a
third order polynomial based on net
generating capacity provided by the user.
The cost of the chimney was estimated with a
power law on total flue gas flow exiting each
absorber, based on CUECost. Separate
power laws were used depending on whether
reheat was included in the design.
The BMs cost was a function of chimney
cost. Chimney cost was estimated with a
power law based on flow rate per absorber.
Temperature at the chimney inlet was
selected in the model at 127 °F, while the
pressure was selected at 4 in. H2O gauge.
The BME cost was estimated as follows:
BME = BARE MODULE^ + CHIMNEY
(6-20)
For a BME with reheat, the cost of chimney
(CHIMNEY 1) was estimated as:
CHIMNEY 1 = 40208 • ACFMY333
(6-21)
For a BME without reheat, the cost of
chimney (CHIMNEY 2) was estimated as:
factors, resulting in an estimate of Total Plant
Cost (TPC). The financial factor D, which
includes the effects of inflation on the cost of
capital and relates to the time required to
construct the FGD equipment, was applied to
the TPC to estimate Total Plant Investment
(TPI). Finally, TCR is the sum of TPI,
inventory cost, and pre-production cost. Pre-
production cost incorporated one-twelfth of
the projected annual O&M expense and 2
percent of the TPI estimate. Detailed
calculations are described below.
Following the EPRI TAG approach, 5 percent
for general facilities, 10 percent for
engineering and home office, 5 percent for
process contingency, and 15 percent for
project contingency were applied. This
model also included a Prime Contractor's Fee
of 3 percent, which is the CUECost algorithm
default. This cost was added to arrive at the
Total Plant Cost (TPC). Using these
CUECost defaults and adding yielded TPC
for the model:
TPC=BM»\l + ^ + ^ + ^
100 100 100
CHIMNEY 2 = 23370 • ACFMl
0.3908
(6-22)
B
Too
• 1 +
c
ioo
(6-24)
The auxiliary cost for Support Equipment
Area was estimated as:
BARE MODULE^ = 0.0003 »MWe3 -1.0677
• MWe2+1993.8 »MWe+1177674 (6-23)
Total Capital Requirement
Once the BM cost had been determined, it
was possible to calculate TCR. The general
TCR determination procedure is illustrated in
Table 6-3. Following the EPRI TAG
methodology, installed BM cost was
multiplied by appropriate contingency
TPC could then be adjusted for financial
factors dependent on the time required to
complete the project. Allowance for Funds
During Construction Factor (FAFDC) and Total
Cash Expended Factor (FTCE) are used to
adjust TPC. FAFDC accounts for interest
during construction and FTCE allows for de-
escalation of cost. CUECost includes time
requirements for various size FGD
installations.
52
-------
Table 6-3. TCR Calculation Method
Cost Component
Symbol / Calculation
Capital Cost
Facilities + Engineering & HOa + Process Contingencies
Project Contingency
Fee
Total Plant Cost (TPC)
Financial Factor
Total Plant Investment
Pre-production Cost + Inventory Capital
Total Capital Requirement
BM = BMF + BMR +BMG + BMW + BME
A = A: + A2 + A3
B
C
TPC = BM * (1 + A) * (1 + B) * (1 + C)
FTCE + FAFDC = D
TPI= TPC *D
E
TCR = TPI + E
" HO = Home Office
Applying the FTCE and FAFDC appropriate to
the unit size results in Total Plant Investment
(TPI):
TPI=TPC.(FTCE+FAPDC)
(6-25)
In regulatory cost determinations, it is usually
preferable to assume constant dollars; e.g. no
inflation. Such analysis should yield a FTCE
equivalent to 1 (no inflation), and an FAFDC
dependent on cost of capital without inflation.
Applying an FAFDC rate of 7.6 percent and
zero inflation results in factors listed in
Table 6-4. Constant dollar factors listed in
Table 6-4 are used in the subsequent model
development.
The Total Capital Requirement (TCR) was
determined by adding pre-production cost and
inventory capital to TPI. CUECost estimates
pre-production cost as a sum of 2 percent of
TPI plus one-twelfth of projected annual
fixed O&M cost plus one-twelfth of projected
annual variable O&M cost adjusted for the
capacity factor, as follows:
TCR = 1.02 »TPI-
VariableO&M
+ CF»12
FixedO&M
I
12
+ INVENTORY
(6-26)
where inventory capital (INVENTORY) is
the cost of reagent required to meet the bulk
storage requirement. A 60-day limestone
inventory was incorporated (limestone cost of
$15/ton was used). CF is plant capacity
factor. CF is defined as a ratio of the average
output to the rated output of a plant on an
annual basis.
Finally, a correction was made to the TCR to
account for the cumulative effect of variables
with minor cost impact (Table 6-2), which
were determined based on the sensitivity
analyses. The CUECost-determined TCR for
baseline conditions shown in Table 6-1 and
for minor effect variables fixed as shown in
Table 6-2 was equal to $205/kW. However,
when minor effect variables were set to
maximize their combined effect on cost, the
resulting value of TCR was $226/kW.
Therefore, TCR was multiplied by the
adjustment factor of 1.1024 (226/205) to
yield the Adjusted TCR.
53
-------
Table 6-4. Financial Factors for FGD Construction, Constant Dollars
Unit Capacity Years to complete AFDC Factor
TCE Factor
MWe < 160
160 < MWe < 400
400 < MWe < 725
725 < MWe < 1300
1300 < MWe < 2000
MWe = 2000
1
2
3
4
5
6
0.0000
0.0380
0.0779
0.1199
0.1640
0.2104
1.0000
1.0000
1.0000
1.0000
1.0000
1.0000
Operation and Maintenance Cost
The O&M cost was calculated next. The
O&M cost includes fixed and variable
components. Fixed O&M cost incorporates:
• operating labor
• maintenance labor and materials
• administration and support labor
Variable O&M cost is composed of:
• reagent
• dibasic acid
• disposal (by-product credit given)
• steam
• electrical energy
Fixed O&M cost components were estimated
as follows. Operating labor (OL) was
estimated by the equation below, using a
power law on the unit's capacity and
estimating the number of workers needed in
combination with an operating labor rate
($30/hr):
OL =
100
(6-27)
Maintenance labor and materials (ML&M)
cost was determined as a percentage (3
percent) of BM cost. Administration and
support (A&S) labor was estimated as a
fraction of maintenance labor and materials
plus operating labor, as given by the
equation:
A & S = 0.3 • (0.4 »ML &M + OL)
(6-28)
Variable O&M cost components were
estimated as a sum of limestone, DBA,
disposal, steam, and electrical energy costs.
Cost of limestone (unit price of limestone at
$15/ton) was:
FRL
2000
(6-29)
where CF is capacity factor.
Cost of dibasic acid (unit price of dibasic acid
at $430/ton):
• 8760 »CF» 430 (6-30)
2000 2000
54
-------
Cost of disposal if the gypsum stacking
method is selected ($6/ton):
CDS = 6 • 8760 • CF • FRSO • 0.95 •
172
64.2000
Cost of disposal if landfilling is selected
($30/ton):
(6-31)
Cnr = 30 • 8760 • CF • FR,n »0.95»
172
--DL
•SO,
64.2000
(6-32)
Cost of disposal was set to zero if wallboard
production was selected. In addition, for this
case a by-product credit ($2/ton) was given as
described below:
CREDIT =
FRSO »0.95»
172
64 • 2000
-•8760»CF»2
(6-33)
Cost of steam (price of steam estimated at
$3.50/1000 Ib):
STEAM = •
TER
855.14.1000
-• 8760 »CF» 3.5
(6-34)
Cost of electrical energy (power consumption
for LSFO estimated at 2.0 percent) was
estimated using the default CUECost power
priceof25mills/kWh:
POWER = 0.02»
• 8760.CF.25
(lOOO»MWe» 0.8231)
1000
(6-35)
As an annual expense, the components of
variable O&M cost were adjusted for the
capacity factor of the unit(s).
Validation
Capital cost predictions of the simplified
LSFO cost model were validated against
reported capital cost for eight recent retrofit
LSFO systems.
LSFO cost estimates derived by the
simplified model described above were
validated against reported costs (CUECost
manual) for eight Phase I plants with retrofit
scrubbers. These eight plants included one
LSIO retrofit, Gibson, and seven LSFO
retrofits of various configurations. Since the
simplified cost model incorporates
generalizations applied to the CUECost
algorithm, it was necessary to validate this
model against these recent retrofits.
Model estimates of TCR and published costs
are presented in Table 6-5 and are further
illustrated in Figure 6-2. These results reflect
that the simplified LSFO cost model on the
average predicts the published capital cost
within 10.5 percent.
In the validation study, a heat rate of 10,500
Btu/kWh and a coal heating value of 11,900
Btu/lb were used for all plants. All of the
Phase I units in Table 6-5 were designed
without reheat. Absorber materials of
construction and the disposal mode for each
unit are shown in Table 6-5. The simplified
model cost was de-escalated to 1994 dollars
to maintain consistency with reported costs.
55
-------
Table 6-5. Model Validation Summary for LSFO FGD (1994 Dollars)
Plant
Petersburg
Cumberland
Conemaugh
Ghent
Bailly
Milliken
Navajo
Absorber Material/
Disposal
Alloy/landfill
RLCS/stacking
RLCS/wallboard
Alloy/stacking
RLCS/wallboard
RLCS/wallboard
Alloy/landfill
Unit Capacity,
MWe
239
1300
1700
511
600
316
750
Absorbers
1
3
5
3
1
1
2
Coal
Wt%S
3.5
4.0
2.8
3.5
4.5
3.2
0.75
Model,
$/kW
400
164
174
213
189
368
226
Reported,
$/kW
317
200
195
215
180
348
236
Deviation,3
percent
+26.2
-18.0
-10.8
-0.1
+5.0
+5.7
-4.2
aDeviation=(Model-Reported)/Reported» 100%
4oU
Ar\r\
4UU
•acn
oOU
onn
ouu
ocn -
ZOU
onn
ZUU
^ Kn
IOU
A r\r\
1UU
Kn
ou
0_
A
A
•
. *
* ' :
-•- Data ($/kW)
-A- Model ($/kW)
0 500 1000 1500 2000
Unit Capacity, MWe
Figure 6-2. Comparison of model predictions with cost data for LSFO.
56
-------
Recently, IPM model predictions63 of TCR's
were published for 2-4 percent sulfur coals.
The comparison of the simplified CUECost
model prediction to the IPM model for 2, 3,
and 4 percent sulfur coals is given in Figure
6-3. As can be seen in Figure 6-3, model
predictions of TCR are not very sensitive to
coal sulfur content for the range of 2 to 4
percent.
State-of-the-art Model
The algorithms developed thus far
incorporated a variety of adjustments to
CUECost algorithms to eliminate variables
that did not have significant impact on cost.
At this point, however, it is helpful to specify
a "state-of-the-art" LSFO system by which to
estimate the cost of possible future retrofits.
It is recognized that alternate design decisions
may be made in the interest of reducing cost
based on site specific conditions or other
engineering features resulting in cost savings
not reflected otherwise.
Therefore, the simplified LSFO cost model
was further adjusted with cost-effective
design decisions to arrive at the LSFO part of
the SUSCM (LSFO SUSCM). This latter
model is expected to provide the budgetary
cost estimates for future LSFO applications.
The assumptions made in arriving at the
LSFO SUSCM are described below.
1. Absorbers serving flue gas from units up
to 900 MWe in capacity are used in the
LSFO SUSCM designs. This is
consistent with the recently reported
information for Units 1 and 2 of Tampa
Electric's Big Bend Station. At this
station, both units were retrofitted with a
single 60-ft diameter 890-MWe
, , 44 64 65
module ' ' .
2. The "state-of-the-art" scrubber is
constructed of rubber-lined carbon steel
or alloy material. Scrubber cost was
assumed to be the average of rubber-lined
carbon steel and alloy materials.
3. The "state-of-the-art" scrubber uses
dibasic acid addition, resulting in modest
capital savings and significant O&M
savings.
4. The "state-of-the-art" scrubber uses
gypsum stacking or wallboard production
as the waste disposal method. Waste
disposal bare module cost was assumed to
be the average of the cost for the two
disposal methods.
5. Sorbent inventory of 30 days.
6. The cost of chimney was assumed to be
the average of chimney cost with and
without reheat.
"State-of-the-art" decisions are shown in
Table 6-6.
Combining the equations developed before
with these "state-of-the-art" design decisions
yields a LSFO SUSCM-derived estimate of
TCR for a "state-of-the-art" FGD unit. TCR
predictions using LSFO SUSCM are shown
in Figure 6-4. These predictions are based on
units with a heat rate of 10,500 Btu/kWh and
a capacity factor of 90 percent. The results
reflect that the capital cost is not sensitive to
coal sulfur content. However, as expected,
capital cost does reflect an economy-of-scale.
It is worth noting that the discontinuities in
capital cost curves reflect the addition of an
absorber as unit capacity changes from less
than 900 MWe to greater than 900 MWe and
from less than 1800 MWe to greater than
1800 MWe. This is because of assumption 1
described above.
57
-------
600
- IPM, 2%S ($/kW)
- IPM, 3% ($/kW)
- IPM, 4% ($/kW)
-Model, 2%S($/kW)
-Model, 3%S($/kW)
-Model, 4%S($/kW)
500 1000 1500
Unit Capacity, MWe
2000
Figure 6-3. Comparison of LSFO cost model to IPM model predictions for 2 to 4 percent sulfur coal.
-------
Table 6-6. "State-of-the-art" LSFO Design Decisions
Parameter Units
Single Absorber Size MWe
Absorber Diameter ft
DBA Addition3
L/G gal/1000 ft3
O2 in Stack %
SO2 Removal %
Flue Gas Temperature from Absorber °F
Flue Gas Velocity into Absorber ft/s
Inventory for Limestone days
Limestone Purity (CaCO3) %
Waste Disposal
Power Requirement %
Flue Gas Reheat3
aYes/No decision only; no addition rate considerations
Value
900
60
Yes
70
8
95
300
14
30
95.3
Average of wallboard or gypsum stacking
2
Average of Yes and No
59
-------
500
400
300
o
o
-SUSCM Model, 2% S
-SUSCM Model, 3% S
- SUSCM Model, 4% S
0 200 400 600 800 1000 1200 1400 1600 1800 2000
Unit Capacity, MWe
Figure 6-4. TCR predictions for 2 to 4 percent sulfur coal by LSFO SUSCM.
-------
For comparison, the published average cost of
24 Phase I units66 was $249/kW (1995
dollars) or $241/kW when de-escalated to
1994 dollars. Significant cost reductions may
be realized by employing "state-of-the-art"
design. For example, the LSFO SUSCM
predicts a LSFO TCR of $211/kW for a 500
MWe system with 4 percent sulfur coal. For
the same conditions, the simplified LSFO
model predicted a TCR of $229/kW.
Setting the LSFO SUSCM parameters to
values representative of conditions at Big
Bend Station resulted in a predicted TCR of
$153/kW (with the TPC of $107 million).
Further, giving the credit for the effect of
high velocity in the absorber, TCR decreases
to $145/kW.
As described earlier, fixed O&M was a
function of the installed BM cost and the unit
capacity (MWe). The LSFO SUSCM
prediction of fixed O&M for a unit with a
heat rate of 10,500 Btu/kWh is shown in
Figure 6-5. The fixed O&M cost is based on
capital cost and, therefore, reflects the same
trends as capital cost. The LSFO SUSCM
prediction for Big Bend Station's fixed O&M
is $6/kW-year.
As can be seen in Figures 6-4 and 6-5, LSFO
SUSCM predictions of TCR and of fixed
O&M are not very sensitive to coal sulfur
content in the range of 2 to 4 percent.
Variable O&M is a function of the sulfur
input and power requirements, adjusted for
capacity factor. The LSFO SUSCM
prediction of variable O&M for a unit with a
10,500 Btu/kWh heat rate and 90 percent
capacity factor is shown in Figure 6-6.
Variable O&M costs on a mills/kWh basis are
constant across the unit capacity range and
increase with fuel sulfur content. The LSFO
SUSCM prediction for Big Bend Station's
variable O&M is 1.37 mills/kWh.
61
-------
18.0
0.0
-SUSCM Model, 2% S
-SUSCM Model, 3% S
- SUSCM Model, 4% S
0 200 400 600 800 1000 1200 1400 1600 1800 2000
Unit Capacity, MWe
Figure 6-5. Fixed O&M predictions for 2 to 4 percent sulfur coal by LSFO SUSCM.
-------
2.0
1
03
o
CG
1.2
1.0
- SUSCM Model, 2% S
- SUSCM Model, 3% S
- SUSCM Model, 4% S
0 200 400 600 800 1000 1200 1400 1600 1800 2000
Unit Capacity, MWe
Figure 6-6. Variable O&M predictions for 2 to 4 percent sulfur coal by LSFO SUSCM.
-------
Lime Spray Drying
Sensitivity Analysis
Sensitivity analyses were performed to
determine variables that have relatively minor
impacts on FGD cost. The objective of these
analyses was to build an order of magnitude
cost estimate model using commonly
available parameters that significantly affect
cost.
For the sensitivity analyses, it was necessary
to identify a baseline LSD system as a point
of reference. A 500-MWe unit with a 10,500
Btu/kWh heat rate burning 1.5 percent sulfur
coal was selected as the baseline unit.
The primary design elements fixed in this
baseline LSD system were the spray dryer
absorber construction materials and stack
construction. RLCS was selected as the
construction material for the baseline unit.
The baseline LSD system uses two absorbers
per CUECost methodology (maximum
absorber size 300 MWe). Other variables
were fixed at CUECost default values for the
baseline LSD, including 90 percent 862
removal efficiency. Thus defined, the
baseline LSD has an annual operating cost of
10.02 mills/kWh.
Results of the sensitivity analyses are
summarized in Table 6-7. Values for the
variables were selected to span realistic
ranges. High and low values of variables
were selected and the corresponding cost was
then determined for each single variable
perturbation. Next, the differences in cost
predictions were calculated between baseline
and high, as well as low, values for each
perturbed variable.
Based on the sensitivity analyses, it appears
that the majority of cost impacts can be
accounted for with capacity, heat rate, coal
sulfur content, and coal heating value.
By fixing variables that have minor impacts
on the cost, the methodology can be reduced
to a function of just a few variables. The
variables that have minor impacts on the cost,
as predicted by CUECost sensitivity analyses,
as well as the respective fixed values, are
shown in Table 6-8. These fixed values are
based on the baseline CUECost case.
Capital Cost
Similarly to LSFO, installed capital cost (BM
cost) for LSD is calculated for each of five
major equipment areas. The estimation
methods used for the five major equipment
areas are described below.
Reagent Feed Area
Reagent Feed Area cost (including receiving,
storing, and slaking) - addition of a linear
component based on the design lime addition
rate (Ib/h) and a power law component based
on fresh lime slurry feed rate (gpm). The
lime addition rate was determined by the
uncontrolled SC>2 emission rate and the coal
sulfur content. Fresh lime slurry feed rate
was calculated for the lime addition rate at 30
percent solids, 1.3 specific gravity, and 90
percent lime purity.
The Reagent Feed Area cost (BMF) was
estimated based on the lime feed rate. Lime
purity has been fixed at 90 percent CaO,
which was used as the default composition in
CUECost. The cost estimate was then
calculated for the coal sulfur content, which
64
-------
Table 6-7. Sensitivity Analysis of LSD Annual Operating Cost (baseline value of 10.02 mills/kWh)
Variable, units Baseline Variable's Variable's Cost for Cost for Low High Value Low Value
High Value Low Value High Value Value of Difference,3 Difference,b
of Variable, Variable, % %
mills/kWh mills/kWh
Capacity, MWe
Heat Rate,
Btu/kWh
Coal Sulfur
Content, %
Coal Heating
Value, Btu/lb
Air Heater
Outlet, °F
SO2 Removal,
Adiabatic
Saturation
Temp, °F
Approach to
Saturation, °F
Recycle Slurry
Solids, %
# of Absorbers
Absorber
Material
500
10,500
1.5
11,922
300
90
127
20
35
2
RLCS
2000
11,000
2.00
14,000
360
95
145
50
50
3
Alloy
100
8,000
1.00
10,500
280
85
110
10
10
N/TC
N/T
4.76
10.29
11.16
9.14
10.10
10.16
10.04
10.05
10.01
10.57
10.69
18.77
8.58
8.86
10.78
9.99
9.88
10.00
10.01
10.10
N/T
N/T
52.5
2.7
11.4
-8.8
0.8
1.4
0.2
0.3
-0.1
5.5
6.7
87.3
-14.4
-11.6
7.6
-0.3
-1.4
-0.2
-0.1
0.8
N/T
N/T
"(Cost for High Value of Variable - 10.02) /10.02. 100%
b(Cost for Low Value of Variable - 10.02) /10.02. 100%
CN/T = not tested
65
-------
Table 6-8. Representative Values for LSD Variables with Minor Cost Impacts
Variable Units Value
Comments
Coal Heating Value Btu/lb
Lime Purity % CaO
SO2 Control Efficiency %
Ambient Pressure in. Hg
Air Heater Outlet Temperature °F
Moisture in the Flue Gas %
Approach to Saturation °F
Adiabatic Saturation Temperature °F
Recycle Slurry Solids %
11,900
90.0
90.0
29.4
300
6.0
14.0
20
127
35
Baseline 11,922
Before control device
After control device
determined the stoichiometric ratio (1.75
taken for 3.43 percent S coal), and by the
maximum feed rate to the FGD system. As
described earlier, the heating value was fixed
at 11,900 Btu/lb in this model. The SO2 flow
rate can be estimated based on the coal sulfur
content, unit capacity [MWe], and heat rate
[Btu/kWh] as follows:
Wt%S*WQQ
HHV
64_
32
MW»HR
(6-36)
where Wt%S is coal sulfur content (wt%)
MWe is LSD size, HR is plant heat rate
(Btu/kWh), and HHV is coal heating value
(Btu/lb).
Once the 862 flow rate is known, the Reagent
Feed Area cost (BMF) may be estimated as
follows:
BMF = 170023 •- + 3764611 |
1000
- (72338.GPM03195)
(6-37)
where FRL is the reagent feed rate:
FRL=FRSO
2
. _. 56 1-0.9
• 1.75* — • -
64 0.9
64
(6-38)
and GPM is slurry flow rate:
56
56 0.3
8.34.
60
(6-39)
SO? Removal Area
SC>2 Removal Area cost (including spray
dryers, tanks, and pumps) - third order
polynomial based on coal sulfur content.
Cost of spray dryers - second order
polynomial based on actual gas flow
rate entering each absorber [cfm]
multiplied by the number of
absorbers. Absorber size was limited
66
-------
in CUECost to treat a maximum of
300 MWe; larger units require
multiple equal size absorbers.
The SC>2 Removal Area cost (BMR) required
estimation of flue gas flows and selection of
absorber materials. Gas flow was calculated
in a manner similar to that used for LSFO
calculations to yield the flow as shown
below:
100
1000 9780 (460 + 295)
ACFM = — — • - • - - '-•-, .
106 60 528 (100-6)
uu
»HR»
0.04 0.209 (P-0.04)
~P~+ P * (0.209-P]
(6-40)
The pressure at the absorber inlet was fixed at
12 in. H2O vacuum (the CUECost default).
Ambient pressure was fixed at the CUECost
default of 29.4 in. Hg. Oxygen at 9.0 percent
was assumed throughout the LSD. The
moisture fraction was assumed to be 6
percent at the spray dryer inlet.
The above assumptions allowed for the
estimation of the SC>2 Removal Area cost
(BMR), as shown below.
BMR =BAREMODULER + SPRAY DRYERS
(6-41)
For the SO2 Removal System with RLCS
construction, the cost of spray dryers
(SPRAY DRYERS 1) was calculated as:
SPRAY DRYERS\ =
-3.57.
ACFM
f • 1000
I A CFM
1+791896
V,, • 1000
(6-42)
• For the SC>2 Removal System with alloy
construction, the cost of spray dryers
(SPRAY DRYERS2) was calculated as:
SPRAY DRYERS2 =
+ 1080990
' •#„
(6-43)
where Na is the number of absorbers.
Auxiliary cost (BARE MODULER) was
calculated as:
BAREMODULER =
1^581877809• Wt%S3 -3653117• Wt%S2\ ^
[+ 693335 »Wt%S + 214198 J* '
+677421<
(6-44)
Flue Gas Handling Area
Flue Gas Handling Area cost (including
ductwork and fans) - linear addition of power
laws based on the actual flue gas flow rate
entering the absorber, exiting the absorber,
exiting the paniculate control device, and
exiting the ID fans.
- Cost of ID fans - power law based on
the flue gas flow rate [cfm] handled
by each fan multiplied by the number
of fans required. The number of fans
required was determined by the total
gas flow rate and the maximum gas
flow rate per fan (1,600,000 cfm).
The Flue Gas Handling Area cost (BMG) was
estimated based on flue gas flow rates at
multiple locations: entering the absorber,
exiting the absorber, exiting the particulate
control device, and exiting the ID fans. The
67
-------
flue gas exiting the absorber was assumed to
be at 17 in. H2O vacuum and 147 °F,
consistent with a 20 °F approach to
saturation. Flue gas exiting the paniculate
control device was assumed to be at 23 in.
H^O vacuum and 147 °F. Flue gas exiting the
fans was assumed to be at 1 in. H^O gauge
positive pressure and at 152 °F. The
CUECost model adjusts flue gas flow rates to
account for water evaporation and acid gas
removal. For flue gas flow estimating
purposes, all flue gas flows after the absorber
inlet had a water content of 14 percent.
The Flue Gas Handling Area cost included
the cost of ID fans. It was estimated using
the flue gas flow rate exiting the paniculate
control device and the number of fans
required. CUECost determines the number of
fans through a series of logical comparisons
based on maximum individual fan capacity at
the specified pressure change across the fan.
The pressure differential across the fans was
fixed at 24 in. H2O.
Based on the assumptions presented above,
the Flue Gas Handling Area cost (BMo) was
estimated as follows:
where ACFM1, ACFM2, and ACFM3 are
flue gas flow rates at the exit from the
absorber, paniculate control device, and ID
fans, respectively. Na is the number of
absorbers.
The cost of ID fans (ID FANS) was
calculated as:
/- \0.6842
ID FANS = 91.24. \ACFM2\ .N
\ Nf } f
where Nf is the number of fans.
Waste/By-product Handling Area
Waste/By-product Handling Area cost
(including disposal and storage) - second
order polynomial based on coal sulfur content
(Wt%S).
Waste Handling Area cost (BMw) was
estimated as a function of coal sulfur content.
Waste included fly ash and was presumed to
be sent to a landfill. BMw was estimated as
follows:
(6-47)
BMW = 205 1841 884
5' -1443163
BMr = BARE MODULE^ + ID FANS
(6 - 45) . Wt%S +1026479
(6 - 48)
The area's auxiliary cost (BARE MODULEG)
was estimated by the following equation:
BAREMODULE =
1721.8.
15338.
4840.4.
0-683
1000 )
1000 )
f ^/-zrj.yi
+ 1326.2. ACFMl]
{ 1000 )
+ 47680.
1000 )
+ 2695.9.
Support Equipment Area
Support Equipment Area cost (including
electrical, water, and air) - second order
polynomial based on the unit capacity
(MWe).
Cost of chimney - power law based on
the flue gas flow rate (ACFM3)
exiting the ID fans.
The Support Equipment Area cost
included the chimney without reheat. The
chimney cost (CHIMNEY) was based on the
68
-------
flue gas flow rate and was estimated as
follows:
CHIMNEY = 23370 • ACFM
(6 - 49)
Support Equipment Area cost
calculated as:
was
BME =- 1.21 l
We +2704.2 • MWe
+ 1354716.2 + CHIMNEY
(6-50)
Adding the BM cost components for the five
major areas yields an estimate for installed
capital cost.
Total Capital Requirement
Once the BM cost had been determined, it
was possible to calculate LSD TCR. Total
Plant Cost (TPC) was calculated in the same
manner as explained before for LSFO in
equation (6-24).
Next, TPC was adjusted for financial factors
dependent on the time required to complete
the project.
As explained before for LSFO, the
adjustment results in Total Plant Investment
(TPI) as described before in equation (6-25).
Since it is usually preferable to assume
constant dollars in regulatory applications, a
constant dollar analysis was done as
explained before in the LSFO section.
Current dollar factors were used for
validation, assuming that the published cost
for TCR was in current dollars. Constant
dollar factors were used in the subsequent
model development.
TCR was determined by adding pre-
production cost and inventory capital to TPI.
CUECost estimates pre-production cost at 2
percent of TPI plus one-twelfth of the
projected annual O&M (fixed plus variable
adjusted for capacity factor) cost. Similar to
considerations for LSFO, a 60 day lime
inventory was incorporated in the model.
The default cost of lime used here was
$50/ton. Substituting the default factors in
TPI and the default cost of lime yields TCR
as described before by equation (6-26).
The CUECost-determined TCR for baseline
conditions shown in Table 6-8 and for minor
effect variables fixed as shown in Table 6-9
was equal to $159/kW. However, when the
minor effect variables were set to yield the
highest cost, the resulting value of TCR was
$165/kW. Therefore, TCR was multiplied by
the adjustment factor of 1.038 (165/159) to
yield the Adjusted TCR.
Operation and Maintenance Cost
The O&M cost was calculated next. The
O&M cost includes fixed and variable
components. The fixed O&M cost
incorporates:
• operating labor
• maintenance labor and materials
• administration and support labor
The variable O&M cost is composed of:
• reagent
• disposal
• fresh water
• energy
Fixed O&M cost components were estimated
as follows. Operating labor (OL) was
estimated by the equation below, using a
power law on unit capacity and estimating
number of workers needed in combination
with an operating labor rate ($30/hr):
69
-------
OL = 18.25 -2.278-MJF •
., ln(MWe)»30»40»52
/ — \ e /
100
(6-51)
The maintenance labor and materials
(ML&M) cost was determined as a
percentage (2 percent) of BM cost.
Administration and support (A&S) labor was
estimated as a fraction of maintenance labor
and materials and of operating labor, as given
by the equation:
A & S = 0.3 • (0.4 • ML & M + OL)
Variable O&M cost components were
estimated as a sum of lime, disposal, fresh
water, and energy costs. The cost of lime
(unit price of lime at $65/ton) was:
(6-52)
CCa0 =
cao
770
• 8760 • CF • 65
(6-53)
where CF is capacity factor.
The cost of disposal ($30/ton) is:
CDL =
8760
2000
• CF • 30
FRSO • —+ MJF« 1000• 0.1 «-^-
" 64 HHV
(6-54)
The cost of energy (energy consumption for
LSD estimated at 0.7 percent) was estimated
using the default CUECost energy price of 25
mills/kWh):
POWER = 0.007 »
(100°
1000
. 8760 . CF . 25 (6-55)
As an annual expense, the components of
variable O&M cost were adjusted for the
capacity factor of the unit.
28
Validation
The 1995 EIA-767 browser database^ on
LSD systems installed in the 1980's has been
used for validation. Six LSD systems were
found in this database with adequate data to
perform validation. However, costs provided
for Stanton 1, East Bend 2, and Craig 3 units
appeared unreasonably low for a FGD system
of this type and were not considered during
validation.
Due to the vintage of these LSD system costs,
it was presumed in modeling that they were
built with RLCS absorbers. Since spray
dryers typically operate between 20 and 30 °F
above the dewpoint, no reheat was assumed
in these designs. Table 6-9 presents
validation data for the LSD model. The
results of validation are also shown in
Table 6-9. Validation of LSD Model
Plant/Unit
H.L. Spurlock/2
Wyodak/1
North Valmy/2
Unit Capacity,
MWe
508
362
267
Coal S, wt %
3.6
0.8
0.5
Number of
absorbers
4
3
3
Reported Cost,
$/kW
189
172
231
Model Cost,
$/kW
222
203
205
Deviation,3 %
17.5
18.0
-11.3
"Deviation = (Model - Reported) / Reported • 100%
70
-------
Figure 6-7. These results reflect that the
simplified LSD cost model on average
predicts the published capital cost within 15.6
percent.
State-of-the-art Model
The algorithms developed thus far
incorporated a variety of adjustments to
CUECost algorithms to eliminate variables
that did not have a significant impact on cost.
At this point, however, it is helpful to specify
a "state-of-the-art" LSD system by which to
measure the cost of possible future retrofits.
It is recognized that alternate design decisions
may be made in the interest of reducing cost
based on site specific conditions, or other
engineering advances, resulting in cost
savings not reflected otherwise.
The model (LSD SUSCM) assumes use of the
minimum number of absorbers possible based
on the maximum size constraint of 275
MWe.67 The "state-of-the-art" LSD used in
the LSD SUSCM incorporates a RLCS
absorber construction, and a 30 day reagent
inventory. "State-of-the-art" LSD design
decisions are shown in Table 6-10.
Combining the equations developed before
with these "state-of-the-art" design decisions
yields a LSD SUSCM-derived estimate of the
TCR for a "state-of-the-art" FGD unit. TCR
predictions using LSD SUSCM are shown in
Figure 6-8.
As described earlier in this chapter, fixed
O&M cost is a function of the installed BM
cost and the unit capacity (MWe). The LSD
SUSCM prediction of fixed O&M cost for a
unit with a heat rate of 10,500 Btu/kWh is
shown in Figure 6-9.
The LSD SUSCM prediction of variable
O&M cost for a unit with a 10,500 Btu/kWh
heat rate and 90 percent capacity factor is
shown in Figure 6-10. Variable O&M costs
on a mills/kWh basis are constant across the
unit capacity range and increase with fuel
sulfur content.
71
-------
250
200
150
o:
o 10°
50
- Data ($/kW)
-Model ($/kW)
0 100 200 300 400 500 600
Unit Capacity, MW(
Figure 6-7. Validation of LSD cost model.
-------
Table 6-10. "State-of-the-art" LSD Design Decisions
Parameter
Units
Value
Single Absorber Size
O2 in Stack
Material of Construction
SO2 Removal
Stoichiometry
Flue Gas Temperature
Lime Inventory
Lime Purity
Lime Cost
Waste Disposal Cost
MWe
days
%
$/ton
$/ton
275
8
RLCSa
90
1.4 for 2% S Coal
300
30
94
50
12
aRLCS = Rubber-lined Carbon Steel
73
-------
400
350
300
250
200
O
50
•SUSCM Model, 1%S
- SUSCM Model, 2%S
500 1000 1500
Unit Capacity, MWe
2000
Figure 6-8. LSD TCR predictions by LSD SUSCM.
-------
12.00
0.00
• SUSCM Model, 1%S
• SUSCM Model, 2%S
500 1000 1500
Unit Capacity, MWe
2000
Figure 6-9. LSD fixed O&M predictions by LSD SUSCM.
-------
ON
2.50
2.00
jfl
^ 1.50
03
o
re
-c
re
1.00
0.50
0.00
H-l—II III I I I I I I I I I I I I I I I » I I I I I » I I I I I I I I
- SUSCM Model, 1%S
-SUSCM Model, 2%S
500 1000 1500
Unit Capacity, MWe
2000
Figure 6-10. LSD variable O&M predictions by LSD SUSCM.
-------
Magnesium-enhanced Lime
General Approach
The approach taken was to estimate the
Magnesium-enhanced Lime (MEL) system
cost, both capital and O&M, based on the
estimation methods previously described for
LSFO and LSD. As described earlier, for
costing purposes, MEL can be considered to
be a combination of LSFO and LSD. The
MEL cost was based on a retrofit presenting a
medium difficulty. The derived algorithm
was then further simplified by making state-
of-the-art design decisions to build a cost
model. TCR was estimated in the same
manner as previously described for LSFO and
LSD.
Capital Cost
The BM was calculated for each of five major
equipment areas, as described before for
LSFO (Reagent Feed, SO2 Removal, Flue
Gas Handling, Waste Handling, and Support
Equipment). Each major equipment area may
have extraordinary items estimated apart from
the rest of the equipment system. The
estimation methods used for the five major
equipment areas were as described below.
The Reagent Feed, SO2 Removal, and Waste
Handling Area cost estimates were explicitly
determined by the SO2 feed rate to the FGD
system. This estimate is determined in
CUECost by the coal sulfur content and coal
use rate with no provision for sulfur retention
in the ash. The higher heating value (HHV)
of the coal was fixed at 11,900 Btu/lb. SO2
feed rate to the FGD system was estimated as
given before in equation (6-1).
Adding the BM cost components from the
five major systems yields an estimate for the
MEL installed capital cost.
Reagent Feed Area
The Reagent Feed Area (BMp) cost
(including receiving, storage, and slaking of
magnesium enhanced lime) was estimated
using the same methodology as the one used
before for the LSD reagent feed area. The
reagent feed ratio remained constant with
respect to coal sulfur content.
The BMp was estimated based on lime feed
rate. Lime purity has been fixed in this
model at 94 percent CaO. Lime addition rate
was fixed in this model at a 1.00 reagent feed
ratio. These parameters allowed the BMp
cost to be estimated as follows:
BMF =
170023 • -^- + 3764611 + 72338 • F«P^0'3195 (6 - 56)
1000
1 GPM
where FRL is reagent feed rate (Ib/hr) and
FGPM is slurry flow rate (gpm).
SO? Removal Area
The SO2 Removal Area (BMR) cost
(including absorber and spray pumps) of the
MEL system is expected to require nominally
the same size and number of tanks as the
LSFO. This system's cost was estimated as a
third order polynomial on SO2 rate to the
scrubber. The cost components were
calculated as follows:
Cost of absorber - Estimated at 90
percent of the cost of LSFO absorbers
to approximate the reduction in height
and elimination of spray headers for
the MEL system. The cost estimate
was based on a power law with the
absorber inlet flow rate to each
absorber multiplied by the number of
absorbers. Separate power laws were
used depending on the absorber
construction materials. Maximum
77
-------
absorber size was limited to 275
MWe; larger units require multiple,
equal size absorbers.
Cost of spray pumps - the same
methodology as previously employed
to estimate LSFO spray pump cost
was applied to MEL (a power law
applied to the slurry flow rate per
absorber per pump multiplied by the
number of pumps). The slurry flow
rate (gpm) was calculated based on
the gas flow rate per absorber at the
exhaust temperature, but at 1 in. H2O
less than the inlet pressure. L/G was
fixed at 40 consistent with the open
tower design and 95 percent SO2
removal.
The BMR cost estimation required calculation
of the flue gas flow through the FGD system.
Tank cost was estimated on the same basis as
the one used for LSFO. Absorber cost was
estimated based on inlet flue gas flow rate
and construction materials. Spray pump cost
was estimated based on gas flow rates
exhausting the absorber.
The flue gas flow rate was calculated in the
same manner as previously explained for
LSFO and LSD. Pressure at the absorber
inlet was fixed at 7 in. H2O gauge, the
CUECost default. Ambient pressure was
fixed at the CUECost default of 29.4 in. Hg.
Temperature of the flue gas entering the
absorber may vary significantly for different
units but is expected to have minimal impact
on TCR, based on the sensitivity analysis for
the LSFO. Absorber inlet temperature was
fixed in the model at 295 °F, resulting from
the 300 °F air heater outlet temperature used
as the default in CUECost. Oxygen at 9.0
percent was assumed at the absorber inlet.
The moisture was assumed to be 6.0 percent
at the absorber inlet.
The cost of spray pumps for the absorbers
was estimated based on the absorber outlet
flow rate and the number of pumps required.
The number of pumps (np) required was
based on the required slurry flow rate per
absorber, and a maximum pump capacity
(43,000 gpm, the same as for LSFO). The
required slurry flow rate was determined by
the L/G, estimated at 40 for 95 percent SO2
removal in an open spray tower. The gas
flow rate was determined at 127 °F and at 6
in. H2O gauge. Moisture content was
estimated at 14 percent.
These approximations allowed estimation of
the BMR cost depending on the material of
construction for the absorber as follows:
BMR =
BAREMODULER + ABSORBER+PUMPS (6 - 57)
• For SO2 Removal Area with alloy
absorber construction:
ABSORBER! =
230064*0. 9 »
1000 )
*N
(6-58)
where ACFM is flue gas flow at the absorber
inlet in cfm and Na is the number of
absorbers.
• For SO2 Removal Area with RLCS
absorber construction:
ABSORBER\=
173978»0.9»
s 0.5575
1000
(6-59)
78
-------
The cost of pumps (PUMPS) was calculated
as follows:
PUMPS = 910.85.
F,
\ 0.5954
GPM
(6-60)
where FGPM is slurry flow rate in gprn and Np
is the number of pumps.
The area auxiliary cost was estimated as
follows:
BAREMODULER =
0.825»
Z7D
r '-Kc-
1000
FRS
- 188.2 •
1000
+ 34809. —2^+1905302
1000
(6-61)
The cost of the fans was estimated by a power
law based on the number of fans required and
the flue gas flow rate. Fans were assumed to
be installed in groups of 2, 4, or 8 with a
maximum fan capacity of 1,600,000 cfm.
The number of fans was based on conditional
tests of the smallest number option (2, 4, or 8)
resulting in an individual fan capacity of less
than 1,600,000 cfm. Inlet pressure for sizing
fans was fixed in the model at 12 in. H^O
vacuum. Temperature at the fan inlet was
fixed in the model at 295 °F.
By fixing these design criteria, the BMo cost
was estimated as follows:
BMG = BAREMODULEG +ID FANS
where area auxiliary cost (BARE
MODULEo ) was:
(6-62)
Flue Gas Handling Area
The Flue Gas Handling Area (BMG) cost
(including ID fans) - MEL was assumed to
have the same flue gas handling requirements
as LSFO. Therefore, cost was estimated with
the same methodology (a polynomial on gas
flow rate entering absorbers, exiting
absorbers, and the number of absorbers).
The BMG cost was based on the number of
absorbers, flow entering absorbers (ACFM),
and flow exiting absorbers (ACFM1).
Pressure of the gas exiting the absorbers was
fixed at 4 in. H2O gauge. The temperature of
the gas exiting the absorbers was fixed at 127
°F, the CUECost default wet bulb
temperature. Flue gas moisture content was
approximated at 14 percent at the absorber
outlet and through the remainder of the FGD
system.
BAREMODULEG =
+ 777.76.^1
1000 i, 1000 J
( ACFM]
+ 238203 - 0.2009 • | - | + 1266.4
I 1000
ACFM]
• I 1 + 420141 + 0.000012.
1000-AT I 1000 I
(ACFM']3
!•
-0.1651.^1 +1288.82.^1
I 1000 I ( 1000
+ 559693
(6-63)
and cost of fans (FANS) was:
FANS = 91.24.
.N,
f
(6-64)
where Nf is the number of fans.
79
-------
Waste/By-product Handling Area
The Waste/By-product Handling Area (BMW)
cost (including thickener and stabilization
equipment) for the MEL waste handling area
was based on LSFO landfill disposal cost (a
third order polynomial on the SC>2 mass flow
rate). If forced oxidation is employed, system
cost would be equivalent to LSFO gypsum
stacking or wallboard by-product options,
also estimated as polynomials based on SO2
mass flow rate. The cost of equipment
components in this model was calculated as
follows:
- Cost of thickener - estimated with the
same method as the one used for
LSFO thickener. Thickener was
estimated as a linear function of waste
solids removal rate.
Cost of stabilization equipment
included a lime bin, ash bin, and small
pugmill to the waste handling system
in addition to components used in the
LSFO algorithm. This additional cost
was included because, for natural
oxidation, waste must be mixed with
lime and fly ash prior to landfilling.
Equipment cost estimates for this
additional equipment were based on a
fraction of Waste Handling Area cost,
including the thickener.
Waste/By-product Handling Area cost (BMW)
was fixed by the disposal option chosen and
by the amount of sludge to be disposed of.
For MEL under natural oxidation, landfill
disposal is the method used by most
installations. This procedure requires similar
equipment as LSFO for landfill disposal but
is sized differently to account for the more
difficult dewatering characteristics of the
MEL waste. The LSFO Waste Handling
Area, excluding the thickener, was presumed
to be dominated by filter cost. This model
assumed 20 percent higher cost based on the
SO2 flow rate compared to LSFO system.
The thickener cost was estimated for LSFO as
a linear function of dry waste disposal rates.
This is consistent with basing cost on the
surface system of the thickener. MEL wastes
from a natural oxidation process require
significantly more surface system per pound
of waste than gypsum wastes due to slower
settling rates. Magnesium salts are expected
to remain in solution and do not affect
settling rates. The amount of paniculate
waste was based on inlet SO2 flow rate,
removal efficiency (fixed at 95 percent), a
reagent feed ratio of 1.05 based on CaO, and
an estimated 5 percent inerts in the lime. For
waste handling cost estimation purposes, all
SO2 removed was assumed to precipitate as
calcium sulfite hemihydrate.
In addition to the waste handling equipment
estimated by CUECost for LSFO, lime and
flyash bins and a pugmill are required. The
total cost of this equipment was estimated at
10 percent of the waste handling system cost,
including the thickener.
BMW cost was estimated as follows:
BMW =
BARE MOD ULEW + THICKENER + D&P
where:
BAREMODULEW =
3
(6-65)
0.325 •
1000
FRV
-168.77.
1000
,
+ 29091. —2- +773243
1000
• 1.25 (6-66)
and
80
-------
THICKENER=
9018.7 »FRSO »0.95»
172
64.2000
-+114562
(6-67)
Bin and pugmill cost (D&P) was 10 percent
of Waste Handling Area.
Support Equipment Area
Support Equipment Area (BME) cost,
including the chimney, was estimated with a
third order polynomial. The cost of the
chimney was estimated based on total gas
flow exiting each absorber.
Support Equipment Area cost (BME) was
estimated as follows:
BME = BARE MODULE + CHIMNEY
(6-68)
where:
BARE MODULE =
0.825»
0.0003 »MWe -1.0667 •MW
I+1993.8.M^e+1177674
2
(6-69)
The chimney cost was estimated with a power
law based on flow rate per absorber in the
same manner as for LSFO. Temperature at
the chimney inlet was fixed in the model at
127 °F, while the pressure was fixed at 4 in.
H2O gauge:
CHIMNEY = 23370 • ACFM\.
0.3908
(6-70)
Total Capital Requirement
Once the BM cost was determined, it was
possible to calculate TCR. First, Total Plant
Cost (TPC) was estimated in the same
manner as previously described for LSFO and
LSD in equation (6-24).
Next, TPC was adjusted for financial factors
depending on the time required to complete
the project. Applying the TCE and FDC
factors appropriate to the unit size (as
explained previously) results in Total Plant
Investment (TPI) as shown before.
In regulatory applications, it is usually
preferable to assume constant dollars; e.g., no
inflation. Therefore, constant dollars were
used in the subsequent model development.
Finally, the Total Capital Requirement (TCR)
was determined in the manner described
earlier in this chapter for LSFO. The cost of
lime of $50/ton was used.68 This lime
typically contains 5 percent MgO.
Substituting the default factors in TPI and the
default cost of lime yielded a TCR prediction.
Operation and Maintenance Cost
O&M cost was calculated next. O&M cost
includes fixed and variable components.
Fixed O&M cost incorporates:
• operating labor
• maintenance labor and materials
• administration and support labor
Variable O&M cost is composed of:
• reagent
• disposal (by-product credit given)
• energy
Fixed O&M cost components were estimated
as follows. Operating labor (OL) cost was
estimated by the equation below, using a
power law on unit capacity and estimating the
number of workers needed in combination
with an operating labor rate ($30/hr):
= 41.69041«MW,
100 • 30 • 40 • 52
(6-71)
81
-------
Maintenance labor and materials (ML&M)
cost was determined as a percentage (3
percent) of BM cost. Administration and
support (A&S) labor was estimated from
maintenance labor and materials and
operating labor as given by the equation
below:
A & S = 0.3 • (0.4 • ML & M + OL)
(6-72)
The variable O&M cost component was
estimated as the sum of lime, disposal, and
energy costs. The cost of lime (unit price of
lime at $50/ton) was:
C
FR
• 8760 • CF • 50
(6-73)
where CF is the capacity factor.
The cost of disposal if gypsum stacking
method is selected ($6/ton) was:
, .0.95. 129
2 64 • 2000
(6-74)
The cost of disposal for landfill ($30/ton)
was:
Cnr = 30 • 8760 • CF • FR,n »0.95»
129
•SO,
64.2000
(6-75)
State-of-the-art Model
At this point, it is helpful to specify a "state-
of-the-art" MEL system by which to measure
the cost of possible future retrofits. Alternate
design decisions may be made in the interest
of reducing cost based on site specific
conditions or other engineering advances
resulting in cost savings not reflected in this
model.
MEL SUSCM will assume use of the
minimum number of absorbers possible,
based on the maximum size constraint of 275
MWe. The "state-of-the-art" MEL scrubber
used in this model incorporates RLCS or
alloy absorber construction and salable
gypsum. "State-of-the-art" MEL design
decisions are shown in Table 6-11.
Combining the equations developed earlier
with these "state-of-the-art" design decisions
yields a model description of a "state-of-the-
art" MEL FGD system.
MEL SUSCM TCR predictions for MEL are
shown in Figure 6-11 for 2, 3, and 4 percent S
coals. These predictions are based on units
with a heat rate of 10,500 Btu/kWh and a
capacity factor of 90 percent. MEL SUSCM
predictions reflect that capital cost is not
sensitive to coal sulfur content.
The cost of energy (energy consumption for
MEL estimated at 1.05 percent) was
estimated using the default CUECost energy
priceof25mills/kWh):
1000
8760.CF.25
(6-76)
As an annual expense, the components of the
variable O&M cost were adjusted for the
capacity factor of the unit(s).
82
-------
Table 6-11. "State-of-the-art" MEL Design Decisions
Parameter
Units
Value
Single Absorber Size
O2 in Stack
Material of Construction
SO2 Removal
L/G
Inventory for Lime
Lime Purity (CaO)
Sorbent Cost
Waste Disposal
Power Requirements
MEL/LSFO Capital Cost Ratio
ID Fans Cost
MWe
gal/1000 ft3
days
$/ton
275
8
Average of RLSC and alloy
98
40
30
94
50
wallboard
1.05
0.80-0.85
2/3ofLSFOIDFansCost
83
-------
The fixed O&M cost prediction is shown in
Figure 6-12. These costs are based on capital
cost and, therefore, reflect the same trends as
capital costs.
Variable O&M cost predictions by MEL
SUSCM are shown in Figure 6-13. Variable
O&M cost on a mills/kWh basis is constant
across the unit capacity range and increases
with fuel sulfur content.
Summary of FGD Cost
The comparison of capital and O&M costs for
three technologies considered here is shown
in Table 6-12. Ranges of costs are given in
1998 constant dollars for a 100 to 1000 MWe
unit. As can be seen in Table 6-12, capital
cost for LSFO used on a small unit (100
MWe) is considerably higher than capital cost
of MEL used on the same size unit. For a
large unit (1000 MWe), capital cost is
comparable for LSFO and for MEL.
Fixed O&M cost is similar for LSFO and
MEL over the entire unit size range
considered. However, variable O&M cost is
lower for LSFO than for MEL, largely due to
the difference in the sorbent cost ($15/ton for
LSFO versus $50/ton for MEL).
Table 6-12. Cost in 1998 Constant Dollars for Selected FGD Technologies
Technology
LSFOb
LSDC
MELd
Capacity Range3
MWe
100 - 1000
100 - 1000
100 - 1000
Capital Cost,
$/kW
542 - 195
363 - 140
384-238
Fixed O&M,
$/kW-Yr
18-7
12-4
16-8
Variable O&M,
mills/kWh
1.80-1.78
2.24-2.24
2.02-2.01
a Unit has a heat rate of 10,500 Btu/kWh and a capacity factor of 90 percent.
b 4.0 percent sulfur coal application, SO2 removal of 95 percent.
c 2.0 percent sulfur coal application, SO2 removal of 90 percent.
d 4.0 percent sulfur coal application, SO2 removal of 96 percent.
84
-------
oo
(Jl
450
400
- SUSCM Model, 2%S
- SUSCM Model, 3%S
- SUSCM Model, 4%S
0 200 400 600 800 1000 1200 1400 1600 1800 2000
Unit Capacity, MWe
Figure 6-11. MEL TCR predictions by MEL SUSCM.
-------
18.00
oo
ON
- SUSCM Model, 2%S
- SUSCM Model, 3%S
- SUSCM Model, 4%S
0.00
0 200 400 600 800 1000 1200 1400 1600 1800 2000
Unit Capacity, MWe
Figure 6-12. MEL fixed O&M predictions by MEL SUSCM.
-------
3.00
2.00
06
o
TO
i.oo
0.00
AAA Jt AAAAAAAAAAAAAAAAAAAAAAAAAAAAAAAAir
-+ » I » » I »»!»»!»! I » I » » I » » I I » I I » I » » I » I I »
• SUSCM Model, 2%S
• SUSCM Model, 3%S
• SUSCM Model, 4%S
0
500 1000 1500
Unit Capacity, MWe
2000
Figure 6-13. MEL variable O&M predictions by MEL SUSCM.
-------
CHAPTER 7
ADDITIONAL BENEFITS
Introduction
The removal of mercury from flue gas by
existing FGD processes could be viewed as
an added benefit of controlling SO2
emissions. Mercury emissions from coal-
fired power generation sources are reported to
be almost 33 percent of the total
anthropogenic emissions in the U.S.69 In
coal-fired power generation, mercury is
volatilized and converted to mercury vapor
(Hg°) in the high temperature regions of
combustion devices. Hg° is transformed into
oxidized mercury (Hg++) as the flue gas cools.
Therefore, the species predominantly present
in flue gas include species of elemental Hg°
and Hg++. It follows that control of both of
these mercury species is necessary to achieve
total mercury emission control.
At present, the control of mercury emissions
from coal-fired boilers is not commercially
practiced in the U.S. The combination of low
mercury concentration and large flue gas
volumes increases the difficulty and cost of
controlling mercury emissions from coal-
fired utility boilers compared to controlling
mercury emissions from municipal waste
combustors.70 However, numerous studies
have been conducted that reported some level
of mercury emission control by the existing
FGD processes. The capability of existing
FGD processes to remove mercury from coal-
fired flue gas is affected by the mercury
species present. Because of mercury's being
the object of particularly strong concern due
to its harmful effects on human health, the
ability of the existing FGD processes to
remove mercury from flue gas is discussed in
more detail in the following sections of this
chapter.
Another added benefit of controlling SC>2
emissions is the effect that decreased
emissions of SC>2 have on the formation of
fine particulate aerosols. July 1997 revisions
to the National Ambient Air Quality
Standards (NAAQS) place emphasis on
particulate matter less than 2.5 |j,m in
aerodynamic diameter (PM^.s).36 These
aerosols are formed in the atmosphere in the
presence of 862 and other gases. Therefore,
an increased scrubber 862 removal
efficiency, leading to lower SC>2 emissions,
may decrease the amount of PM2.5. Source
emissions characterization is required to
understand the fate of aerosol precursors
(such as 802) in the particle formation
process in the atmosphere.71 While PM2.5 can
be produced directly by a variety of sources,
it can also be produced by atmospheric
reactions in the presence of SC>2, NOx, and
VOCs emitted from stationary sources.72
SC>2 is a precursor for sulfuric acid and
sulfate secondary PM2.5 particles. Sulfate
accounts for approximately 47 percent of
PM2.5 in the eastern United States.73 One
strategy to control PM2.5 emissions from
stationary coal-burning sources is to upgrade
the existing particulate control device. The
alternative route may be to control PM2.5
precursors, most notably 862. In this latter
-------
case, modern, state-of-the-art 862 scrubbers,
designed primarily for high efficiency control
of SC>2, could provide an additional benefit by
controlling PM2.5 precursors.
Once-through Wet FGD
A wide range of total mercury removal
efficiency has been reported for once-through
wet FGD applications on bituminous-coal-
fired power generation units. Existing
conventional wet scrubbers can remove
water-soluble Hg++ compounds (e.g.,
mercuric chloride) from flue gas. However, a
major part of Hg°, which is insoluble in water
and the most volatile of the trace metal
species, may pass through wet FGD and
particulate matter control devices.74
Therefore, should the control of mercury
emissions be desired beyond the inherent
control by once-through wet FGD, Hg° would
need to be adsorbed by the sorbent or
converted by reagents or catalysts to a soluble
form of mercury that could be collected by a
wet FGD process.
A mercury measurement program conducted
on six full-scale coal-fired boilers equipped
with ESP and limestone or lime FGD
processes demonstrated an average total
mercury removal across the wet FGD system
of 54 percent (ranging from 45 to 67
percent).75 The ESP inlet and stack flue gas
speciation data indicated 80 to 95 percent
removal of Hg++ across the ESP and wet FGD
system combination. This test program
showed that mercury was also removed by
the fly ash particles (and occasionally bottom
ash). The total mercury removal (defined as
the difference between the mercury input
based on coal firing rate and coal mercury
concentration and mercury stack emissions)
ranged from 59 to 75 percent and averaged 67
percent. It should be noted that the results of
this study were obtained during routine wet
FGD operations and no adjustments were
made to maximize mercury removal.
The statistical analysis of results in the above
program showed a significant correlation
between oxidized mercury removal and
scrubber slurry pH, with a higher pH
resulting in higher mercury removal. Among
coal parameters (all coals included in the
program were mid-chlorine coals), the coal
oxygen concentration showed a strong
negative correlation with oxidized mercury
removal. A weaker correlation was identified
between nitrogen and ash content of coal and
total mercury removal.
Another study on mercury capture by wet
FGD revealed that it could be affected both
by the scrubber design (open spray versus
tray tower) and operational parameters, such
as pH and L/G of the absorber.76 Mercury
emissions from systems equipped with wet
FGD decreased with increasing L/G in the
range from approximately 30 up to
approximately 130 (gal/1,000 ft3). The
decrease of mercury emissions was due to the
decrease of oxidized mercury emissions.
Elemental mercury emissions following the
scrubber remained fairly consistent over the
tested range of operating conditions and the
outlet elemental mercury concentration was
approximately the same as the inlet one.
Operation of the scrubber with the gas flow
distribution tray enhanced mercury removal
over the above L/G range. For example, at
the L/G of 100 (gal/1,000 ft3) the mercury
emissions for a system with tray scrubber
were on the average 38 percent lower than
these measured for a system with the open
spray scrubber. Pilot-scale tests have
demonstrated the potential for removing
approximately 85 percent of the total mercury
emissions using a wet limestone process, with
89
-------
a scrubber configured as a tray tower and
operated at an L/G of approximately 70
(gal/1,000 ft3).74
However, some sampling efforts have
indicated an apparent re-emission of Hg° at
the outlet of wet FGD systems.77 The results
of triplicate measurements revealed from 7.1
to 38.5 percent increases of Hg° concentration
across a wet FGD system operating on flue
gas with inlet concentrations of from 2400 to
2900 ppm SO2.
As discussed before, should the control of
mercury emissions by a wet FGD process
alone be desired beyond the inherent control
by the existing once-through wet FGD, Hg°
would need to be converted by reagents or
catalysts to a soluble form of mercury that
could be collected.
Therefore, bench-scale research and pilot
studies are currently underway to more fully
understand the oxidation of Hg° upstream of
and subsequent to removal in FGD systems.78
The study concentrates on determining
whether the catalyst remains active for
mercury oxidation after an extended exposure
to utility flue gas. One of the findings from
the bench-scale phase of this study is that
HC1 may be participating in the oxidation
mechanism of elemental mercury.
Another route pursued on a bench scale is to
find liquid additives that, once atomized into
the flue gas, would be capable of oxidizing
Hg°. 79 Commercial solutions of chloric acid
and sodium chlorate were capable of
transferring 10 percent of Hg° into solution.
Additionally, approximately 80 percent of the
nitric oxide was removed. Further pilot-scale
evaluations continue to examine mercury
speciation and to develop control options.
Dry FGD
Similarly to wet FGD performance discussed
above, a wide range of 55 to 96 percent
reduction in mercury emissions has been
shown with spray dryers installed on full-
scale, bituminous-coal-fired boilers.74 A
significantly lower reduction of 6 to 23
percent was reported for some
subbituminous-coal-fired boilers. It is
thought that the higher mercury removal
efficiencies seen on bituminous-coal-fired
boilers are related to the higher coal chlorine
concentration in these coals, compared to
subbituminous coals.74 Pilot-scale tests with
the spray drying process have demonstrated a
64 percent total mercury emission reduction
across the spray dryer with 68 percent of total
inlet mercury being oxidized mercury.74
Another dry FGD process that is capable of
additionally removing mercury is the CFB.
Recently presented results of the pilot-scale
testing of a CFB process for mercury
adsorption80 indicated approximately 50
percent of the total mercury removal by
hydrated lime alone and up to 80 percent
removal with the supplemental injection of
iodine-impregnated activated carbon into the
CFB.81 Only total mercury removal has been
tested, mercury speciation in the flue gas was
not reported, and there were no attempts
made to speciate mercury.
The duct injection process may also be used
to control mercury emissions. If, in this
process, a sorbent appropriate for mercury
capture, such as activated carbon or zeolite82,
is co-injected along with the sorbent for SO2
capture, then emissions of SO2 and mercury
may be reduced. In this context, research on
modified hydrated lime sorbents has been
reported.83 However, the duct injection
process has been used sparingly and is
90
-------
considered, at present, to be a niche
application.
In summary, the amount of mercury removed
in an unmodified FGD system is believed to
be a function of mercury speciation. Wet
FGD systems may be able to remove
approximately half of the total mercury from
the flue gas, depending on the coal fired.
Similarly, spray dryers have been found to be
able to remove between 6 and 96 percent of
total mercury, depending on the type of coal
fired. Currently, bench- and pilot-scale
research is underway to more fully
understand mercury speciation and develop
enhanced FGD or stand-alone mercury
control options.
91
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Air Pollutant Control Symposium: The Mega
Symposium: 862 Control Technologies and
Continuous Emission Monitors, EPRI, Palo
Alto, CA, U.S. Department of Energy,
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Protection Agency, Air Pollution Prevention
and Control Division, Research Triangle
Park, NC, 1997. TR-108683-V2.
52 Milobowski, M.G., "WFGD System
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Combined Utility Air Pollutant Control
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Control Technologies and Continuous
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U.S. Department of Energy, Pittsburgh, PA,
and U.S. Environmental Protection Agency,
Air Pollution Prevention and Control
Division, Research Triangle Park, NC, 1997.
TR-108683-V2.
53 Dille, E., K. Frizzell, and W. Shim, "Use of
an Electrochemical Technique for Controlling
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DOE-EPA Combined Utility Air Pollution
Control Symposium: The Mega Symposium:
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CA, U.S. Department of Energy, Pittsburgh,
PA, and U.S. Environmental Protection
Agency, Air Pollution Prevention and Control
Division, Research Triangle Park, NC, 1999.
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54 Chang, J.C.S., and T.G Brna,
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55 Brogren, C., and J.S. Klingspor, "Impact of
Limestone Grind on WFGD Performance,"
EPRI-DOE-EPA Combined Utility Air
Pollutant Control Symposium: The Mega
Symposium: SO2 Control Technologies and
Continuous Emission Monitors, EPRI, Palo
Alto, CA, U.S. Department of Energy,
Pittsburgh, PA, and U.S. Environmental
Protection Agency, Air Pollution Prevention
and Control Division, Research Triangle
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56 Sarkus, T., P. Styf, and S. Vymazal, "Two
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57 Brown, G.N., K.E. Janssen, and P. A.
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for the First Ammonia Scrubbing System,"
EPRI-DOE-EPA Combined Utility Air
Pollutant Control Symposium: The Mega
Symposium: SO2 Control Technologies and
Continuous Emission Monitors, EPRI, Palo
Alto, CA, U.S. Department of Energy,
Pittsburgh, PA, and U.S. Environmental
Protection Agency, Air Pollution Prevention
and Control Division, Research Triangle
Park, NC, 1997. TR-108683-V2.
58 Sedman, C. B., "Controlling Emissions
from Fuel and Waste Combustion," Chemical
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59 Ellison, W., "Worldwide Progress in
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CA, U.S. Department of Energy, Pittsburgh,
PA, and U.S. Environmental Protection
Agency, Air Pollution Prevention and Control
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60 Borio, D., P. Rader, and M. Walters,
"Ammonia Scrubbing: Creating Value from
SO2 Compliance," EPRI-DOE-EPA
Combined Utility Air Pollution Control
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CA, U.S. Department of Energy, Pittsburgh,
PA, and U.S. Environmental Protection
Agency, Air Pollution Prevention and Control
Division, Research Triangle Park, NC, 1999.
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61 Walsh, M.A., "New Marsulex Technology
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EPRI-DOE-EPA Combined Utility Air
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Symposium: Volume 1: SO2 Controls, EPRI,
Palo Alto, CA, U.S. Department of Energy,
Pittsburgh, PA, and U.S. Environmental
Protection Agency, Air Pollution Prevention
and Control Division, Research Triangle
Park, NC, 1999. TR-113187-V1.
62 Keith, R., R. Blagg, C. Burklin, B.
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63 "Analyzing Electric Power Generation
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64 Schimmoller, B.K., "Balancing
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65 Jones, C.,"Meeting Compliance and Profit
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66 Ellerman, A.D., R. Schmalensee, P.L.
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Center for Energy and Environmental Policy
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67 Personal communication from Bjarne
Rasmusseu of Niro A/S, Denmark, to
Wojciech Jozewicz of ARCADIS, USA, July
31,2000.
68 Personal communication from Manny Babu
of Carmeuse Chimie Minerale, Pittsburgh,
PA, to Ravi Srivastava, May 3, 2000.
69 Brenner, R., "Framework for the Future,"
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70 Srivastava, R.K., C.B. Sedman, and J.D.
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71 "Research Priorities for Airborne
Particulate Matter II: Evaluating Research
Progress and Updating the Portfolio," Samet,
J., Committee Chair; National Research
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72 Walker, K.D., K. Fritsky, S.J. Miller, G.L.
Schelkoph, and G.E. Durham, "The Future of
Fine Particulate Matter and Hazardous Air
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73 "Controlling Particulate Matter Under the
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74 Nolan, P.S., G.A. Farthing, D.M.
Yurchison, and MJ. Holmes, "Development
of Mercury Emissions Control Technologies
for the Power Industry," EPRI-DOE-EPA
Combined Utility Air Pollution Control
Symposium: The Mega Symposium:
Volume 3: Particulate Controls, Air Toxics
Control, and Poster Sessions, EPRI, Palo
Alto, CA, U.S. Department of Energy,
Pittsburgh, PA, and U.S. Environmental
Protection Agency, Air Pollution Prevention
and Control Division, Research Triangle
Park, NC, 1999. TR-113187-V3.
75 DeVito, M.S., and W.A. Rosenhoover,
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76 Redinger, K.E., A.P. Evans, R.T. Bailey,
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inFGD Systems," EPRI-DOE-EPA
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Control Technologies and Continuous
Emission Monitors, EPRI, Palo Alto, CA,
U.S. Department of Energy, Pittsburgh, PA,
and U.S. Environmental Protection Agency,
Air Pollution Prevention and Control
Division, Research Triangle Park, NC, 1997.
TR-108683-V2.
77 Brown, T.D., D.N. Smith, R.A. Hargis, Jr.,
and W.J. O'Dowd, "Mercury Measurement
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78 Richardson, C.F., G.M. Blythe, T.R. Carey,
R.G. Rhudy, and T.D. Brown, "Enhanced
Control of Mercury by WetFGD Systems,"
EPRI-DOE-EPA Combined Utility Air
Pollution Control Symposium: The Mega
Symposium: Volume 3: Particulate Controls,
Air Toxics Control, and Poster Sessions,
EPRI, Palo Alto, CA, U.S. Department of
Energy, Pittsburgh, PA, and U.S.
Environmental Protection Agency, Air
Pollution Prevention and Control Division,
Research Triangle Park, NC, 1999.
TR-113187-V3.
79 Livengood, C.D., and M.H. Mendelsohn,
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and Nitric Oxide," EPRI-DOE-EPA
Combined Utility Air Pollution Control
Symposium: The Mega Symposium:
Volume 3: Particulate Controls, Air Toxics
Control, and Poster Sessions, EPRI, Palo
Alto, CA, U.S. Department of Energy,
Pittsburgh, PA, and U.S. Environmental
Protection Agency, Air Pollution Prevention
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and Control Division, Research Triangle
Park, NC, 1999. TR-113187-V3.
80 Helfritch, D.I, P.L. Feldman, and E.G.
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for Mercury Adsorption and Particle
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1999.
81 Helfritch, D.I, P.L. Feldman, and E.G.
Waugh, "The Pilot Scale Testing of a
Circulating Fluid Bed for Mercury
Adsorption and Particle Agglomeration,"
EPRI-DOE-EPA Combined Utility Air
Pollution Control Symposium: The Mega
Symposium, Volume 3: Particulate Controls,
Air Toxics Control, and Poster Sessions,
EPRI, Palo Alto, CA, U.S. Department of
Energy, Pittsburgh, PA, and U.S.
Environmental Protection Agency, Air
Pollution Prevention and Control Division,
Research Triangle Park, NC, August 1999.
82 Morency, J.R., T. Panagiotou, and R.F.
Lobo, "Control of Mercury Emissions in
Utility Power Plants," EPRI-DOE-EPA
Combined Utility Air Pollution Control
Symposium: The Mega Symposium, Volume
3: Particulate Controls, Air Toxics Control,
and Poster Sessions, EPRI, Palo Alto, CA,
U.S. Department of Energy, Pittsburgh, PA,
and U.S. Environmental Protection Agency,
Air Pollution Prevention and Control
Division, Research Triangle Park, NC,
August 1999.
83 Ghorishi, S.B., C.F. Singer, and C.B.
Sedman, "Preparation and Evaluation of
Modified Lime and Silica-lime Sorbents for
Mercury Vapor Emissions Control," EPRI-
DOE-EPA Combined Utility Air Pollution
Control Symposium: The Mega Symposium,
Volume 3: Particulate Controls, Air Toxics
Control, and Poster Sessions, EPRI, Palo
Alto, CA, U.S. Department of Energy,
Pittsburgh, PA, and U.S. Environmental
Protection Agency, Air Pollution Prevention
and Control Division, Research Triangle
Park, NC, August 1999.
100
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