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, D.C. 20460
ii

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.
iii

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
iv

Abstract

Sulfur dioxide (SO2) scrubbers may be used by electricity generating units to meet the 

requirements of Phase II of the Acid Rain SO2 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.
v

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 Injection ....................................................................................... 14

Furnace Sorbent Injection .................................................................................. 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
vi

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
vii

Introduction................................................................................................................... 88

Once-through Wet FGD ................................................................................................ 89

Dry FGD....................................................................................................................... 90

References .......................................................................................................................... 92
viii

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 SO2 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 of LSD 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
ix

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 SO2 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
x

List of Symbols

Symbol Meaning Unit

A&S Administration and support cost dollars

ABSORBER Absorber cost dollars

ABSORBER 1 RLCS absorber cost dollars

ABSORBER 2 Alloy absorber cost dollars

ACFM Flue gas flow into absorber cfm

ACFM1 Flue gas flow out of absorber cfm

ACFM2 Flue gas flow out of ID fans cfm

ACFM3 Flue gas flow out of particulate control device cfm

BARE MODULEE Support equipment area auxiliary cost dollars

BARE MODULEG Flue gas handling area auxiliary cost dollars

BARE MODULER SO2 removal area auxiliary cost dollars

BARE MODULEW Waste handling area auxiliary cost dollars

BM Capital cost for FGD system dollars

BME Capital cost component for support equipment area dollars

BMF Capital cost component for reagent feed area dollars

BMG Capital cost component for waste handling area dollars

BMR Capital cost component for waste handling area dollars

BMW Capital cost component for waste handling area dollars

CB&H Cost of ball mill and hydroclones dollars

CCaCO3 Cost of limestone dollars

CCaO Cost of lime dollars

CDBA Cost of DBA tank dollars

CDL Cost of disposal with landfilling dollars

CDS Cost of disposal with gypsum stacking dollars

CF Capacity factor %

CHIMNEY Cost of chimney dollars

CREDIT By-product credit Dollars
xi

Symbol Meaning Unit

FAFDC Allowance for funds during construction factor -------

Fd F-factor scf/106 Btu

FGPM Slurry flow rate gal/min

FixedO&M Fixed operation and maintenance cost dollars

FRL Reagent feed rate lb/hr

FRSO2 SO2 feed rate to the FDG System lb/hr

FTCE Total cash expended factor -------

HHV Coal heating value Btu/lb

HR Plant heat rate Btu/kWh

ID FANS Cost of ID fans dollars

L/G Liquid-to-gas ratio gal/1000 ft3

ML&M Maintenance, labor, and materials cost dollars

Na Number of absorbers -------

Nf Number of fans -------

Np Number of pumps -------

OL Operating labor cost dollars

P Percent oxygen in the stack %

POWER Cost of electrical energy dollars

PUMP Cost of pump dollars

PUMPS Cost of pumps dollars

SPRAY DRYERS Cost of spray dryers dollars

SPRAY DRYERS1 Cost of RLCS spray dryers dollars

SPRAY DRYERS2 Cost of alloy spray dryers dollars

STEAM Cost of steam dollars

TER Thermal energy required Btu

THICKENER Cost of thickener dollars

VariableO&M Variable operation and maintenance cost dollars

Wt%S Coal sulfur content Wt%
xii

List of Acronyms and Abbreviations

CAAA Clean Air Act Amendments of 1990

CFB Circulating Fluidized Bed

CUECost Coal Utility Environmental Cost Workbook

DBA Dibasic Acid

DSI Duct Sorbent Injection Process

EPA United States Environmental Protection Agency

ESP Electrostatic Precipitator

FGD Flue Gas Desulfurization

FSI Furnace Sorbent Injection Process

ID Inside Diameter

L/G Liquid-to-gas Ratio

LSD Lime Spray Drying Process

LSFO Limestone Forced Oxidation Process

LSIO Limestone Inhibited Oxidation Process

MEL Magnesium-Enhanced Lime Process

MWe Unit Electrical Generating Capacity

NAAQS National Ambient Air Quality Standards

O&M Operation and Maintenance

PM2.5 Particulate Matter Less than 2.5 :m (Aerodynamic Diameter)

RLCS Rubber-Lined Carbon Steel

SUSCM State-of-the-art Utility Scrubber Cost Model

TCR Total Capital Requirement

TPC Total Plant Cost

TPI Total Plant Investment

WESP Wet Electrostatic Precipitator
xiii

Conversion Table – English Units to SI Units

To Obtain From Multiply by

m ft 0.3048

m

2

ft29.29 · 10-2

m

3

ft32.83 · 10-2

°C °F 5/9 (°F – 32) 

kg lb 0.454

J/kg Btu/lb 1.33 · 10-4

m

3

/s cfm 4.72 · 10-4

m

3

/s gpm 6.31 · 10-5

J/kWh Btu/kWh 1055.056

mills $ 0.001

kg/m2in. Hg 345.31
1

CHAPTER 1

INTRODUCTION

Combustion of sulfur-containing fuels, such 

as coal and oil, results in sulfur dioxide (SO2) 

formation. SO2 emissions are known to cause 

detrimental impacts on human health and the 

environment. The major health concerns 

associated with exposure to high 

concentrations of SO2 include breathing 

difficulties, respiratory illness, and 

aggravation of existing cardiovascular 

disease. In addition to the health impacts, 

SO2 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, SO2 and its particulate matter 

derivatives contribute visibility degradation.

Electric power generating units account for 

the majority of SO2 emissions in the U.S. In 

1998, these units contributed 64 percent of 

the national SO2 emissions.1 To mitigate SO2

emissions from electric power generating 

units, the Acid Rain SO2 Reduction Program2

was established under Title IV of the Clean 

Air Act Amendments of 1990 (CAAA). This 

two-phase program was designed to reduce 

SO2 emissions from the power generating 

industry. 

Phase I of the Acid Rain SO2 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 SO2 (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 SO2 Reduction 

Program began on January 1, 2000. The 

nationwide cap for SO2 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 SO2

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, 
2

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 

SO2 control strategies at sources, (2) persons 

involved in developing SO2 and other 

regulations, (3) State regulatory authorities 

implementing SO2 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.
3

CHAPTER 2

FGD TECHNOLOGY

Introduction

Various technologies exist that have been 

designed to remove SO2 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 oncethrough and regenerable, depending on how 

sorbent is treated after it has sorbed SO2.

4

 In 

once-through technologies, the SO2 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 SO2 is released from the 

sorbent during the regeneration step and may 

be further processed to yield sulfuric acid, 

elemental sulfur, or liquid SO2. The 

regenerated sorbent is recycled in the SO2

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, oncethrough 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 CoalPower3 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 majority 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 
4

Limestone Forced Oxidation

Limestone Inhibited Oxidation

Lime

Magnesium-Enhanced Lime

Seawater

Wet

Lime Spray Drying

Duct Sorbent Injection

Furnace Sorbent Injection

Circulating Fluidized Bed

Dry

Once-through

Sodium Sulfite

Magnesium Oxide

Sodium Carbonate

Amine

Wet

Activated Carbon

Dry

Regenerable

Flue Gas Desulfurization

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 

counterflow vertically oriented spray tower. 

A diverse group of wet FGD processes have 

evolved to take advantage of particular 

properties of various sorbents and/or byproducts. 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. 

SO2-containing 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, SO2 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 SO2.
FGD TECHNOLOGY

5

Figure 2-2. Baseline wet FGD system.

Limestone Water

Crushing

Station

Slurry

Preparation

Tank

Flue Gas

In

Flue Gas

Out

Reaction

Tank

Dewatering

Disposal

Absorber

Slurry Bleed

Chimney

Process Water

5
6

The overall reactions in the absorber and in 

the reaction tank can be summarized by:

 (2 -1) 2 2 2

1

3

2 2

1

2 3

CaSO H O CO

SO CaCO H O

• +

+ + →

and

2 (2 - 2)

2

4 2 2

2 2 3 2

1

2

CaSO H O CO

SO O CaCO H O

• +

+ + + →

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 SO2

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 SO2

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, CaSO4 is easier to 

dewater than CaSO3.

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 SO2 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 
7

have a perforated tray added for the 

improvement of SO2 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 SO2 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 SO2 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 SO2 removal in the absorber. The 

amount of available alkalinity for the reaction 

with SO2 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 

O2 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 SO2

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 SO2

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 SO2

(increasing sulfur content of fuel) will 

decrease SO2 removal efficiency by a wet 

absorber. This decreased efficiency is 

observed because increasing SO2

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, 
8

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 wet FGD 

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 byproduct 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
9

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 unconfined 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 sitespecific 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 
10

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 magnesiumenhanced 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.
11

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 SO2. 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 countercurrent 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 
12

sorbent injection process (FSI), and 

circulating fluidized bed process (CFB)]. 

The LSD and the CFB require dedicated 

absorber vessels for sorbent to react with 

SO2, 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 SO2 removals 

below those attainable with wet once-through 

FGD.

Lime Spray Drying

LSD for the control of SO2 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 

particulate 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 SO2 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 SO2 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:

( ) (2 - 3) 2 2

1

2 2

1 Ca OH 2 + SO2 → CaSO3

• H O + H O

( ) 2 (2 - 4) Ca OH 2 + SO3 + H2O → CaSO4• H2O

 (2 - 5) 2 2 4

1 CaSO3 + O → CaSO

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 

SO2 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 SO2

afforded by the recycle promotes increased 

sorbent utilization.
13

Figure 2-3. Lime spray dryer FGD system.

Ball Mill

Slaker

Quicklime

Water

Flue Gas

Recycle Loop

Flue Gas

Out

Disposal

Control

Device

Particulate Feed Tank

Recycle Solids Slurry Tank

Spray

Dryer

Chimney

13
14

Mass transfer during a spray drying process 

occurs in two discrete phases: moist and 

dry.17 During the moist phase, SO2 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, SO2 diffuses 

through the products of the lime and SO2

reaction and causes a gas-solid reaction with 

the unreacted core of lime particle.

Studies indicated that a majority of SO2

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 SO2. 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 SO2 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 SO2 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:

( ) (2 - 6) Ca OH 2 → CaO + H2O

SO2 in the flue gas reacts with the nascent 

CaO as given below: 

 (2 - 7) 2 2 4

1 CaO + SO2 + O → CaSO

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.
15

Figure 2-4. Schematic of DSI.

Control

Device

Particulate

Disposal

Boiler

Hydrated Lime

Silo

Water

Flue Gas

Out

Chimney

15
16

Figure 2-5. Schematic of FSI. 

Control

Device

Particulate

Disposal

Boiler

Flue Gas

Out

Chimney

Humidifier

Reactivation

Reactor

Sorbent

Injection

16
17

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 SO2. 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 SO2 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 

SO2. The liquor is subsequently regenerated 

in evaporators that crystallize sodium sulfite. 

Concentrated SO2 is suitable for sulfuric acid 

production.

Magnesium Oxide

In the magnesium oxide process, SO2 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 SO2

absorption in a scrubber. The absorbed 

product is dried and calcined in a kiln to 

regenerate magnesium oxide. SO2 captured 

during calcination is suitable for sulfuric acid 

production.
18

Figure 2-6. Schematic of CFB.

Control

Device

Particulate

Disposal

Boiler

Flue Gas

Out

Chimney

Water

Circulating Fresh Sorbent

Fluidized Bed

Reactor

Precollector

18
19

Sodium Carbonate 

In this process, SO2 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 

SO2 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 SO2 on 

a moving bed of granular activated carbon. 

Activated carbon is thermally regenerated to 

produce a concentrated SO2 stream. SO2 may 

then be treated by conventional technologies 

to produce sulfuric acid.
20

CHAPTER 3

TECHNOLOGY APPLICATIONS

Introduction

As discussed before, FGD technology 

applications were reviewed based on the 

information in CoalPower3 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 FGDcontrolled 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.
21

Figure 3-1. Historical application of FGD technology in the United States.

0

20000

40000

60000

80000

100000

120000

1970 1972 1974 1976 1978 1980 1982 1984 1986 1988 1990 1992 1994 1996 1998

YEAR INSTALLED

e

REGENERABLE

DRY

WET

21
22

Figure 3-2. Historical application of FGD technology throughout the world.

0

50000

100000

150000

200000

250000

1970 1972 1974 1976 1978 1980 1982 1984 1986 1988 1990 1992 1994 1996 1998

YEAR INSTALLED

e

REGENERABLE

DRY

WET

22
23

Figure 3-3. Wet FGD technology application in the United States.

0

10000

20000

30000

40000

50000

60000

70000

80000

90000

1970 1972 1974 1976 1978 1980 1982 1984 1986 1988 1990 1992 1994 1996 1998

YEAR INSTALLED

MWe

Mag LIME

LIME

SODIUM CARBONATE

LIMESTONE

23
24

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-fuelfired electric power plants through 1998. 

FGD systems were installed to control SO2

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 byproduct 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.
25

Figure 3-4. Dry FGD technology application in the United States.

0

2000

4000

6000

8000

10000

12000

14000

16000

1974 1976 1978 1980 1982 1984 1986 1988 1990 1992 1994 1996 1998

YEAR INSTALLED

MWe

CIRCULATING FLUID BED

FURNACE INJECTION

SPRAY DRYING

DUCT SORBENT INJECTION

25
26

Figure 3-5. Regenerable FGD technology application in the United States.

0

500

1000

1500

2000

2500

3000

1978 1980 1982 1984 1986 1988 1990 1992 1994 1996 1998

YEAR

INSTALLED

MWe

MgO

SODIUM SULFITE

OTHER

26
27

Figure 3-6. Percent shares (capacity) of the three FGD technologies installed.

Wet

Dry

Regenerable

U.S.

82.9%

14.2%

2.9%

Abroad

89.8%

8.3% 1.9%

Worldwide

86.8%

10.9%

2.3%

27
28

Table 3-1. Coal-fired Electrical Generation Capacity (MWe) Equipped with FGD Technology (1998)

Technology United States Abroad World Total

Wet 82,092 114,800 196,892

Dry 14,081 10,654 24,735

Regenerable 2,798 2,394 5,192

Total FGD 98,971 127,848 226,819

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.
29

Table 3-2. Total Capacity (MWe) Equipped with Wet FGD Technology (1998)

Process United States Abroad World Total

Limestone 56,560 106,939 163,499

Lime 14,237 4,338 18,575

MEL 8,464 50 8,514

Sodium Carbonate 2,756 - 2,756

Seawater 75 1,050 1,125

Regenerable (other) - 2,423 2,423

Total Wet FGD 82,092 114,800 196,892

Table 3-3. Total Capacity (MWe) Equipped with Dry FGD Technology (1998)

Process United States Abroad World Total

Spray Drying 11,315 6,904 18,219

Dry Sorbent Injection 2,400 1,125 3,525

CFB 80 517 597

FSI 286 2,108 2,394

Total Dry FGD 14,081 10,654 24,735
30

Figure 3-7. Comparison of limestone and non-limestone wet FGD applications.

Limestone

Non-Limestone

U.S.

82,092 MWe

68.9%

31.1%

Abroad

114,800 MWe

93.2%

6.8%

World

196,892 MWe

83.0%

17.0%

30
31

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 United States Abroad World Total

Wet 174 348 522

Dry 54 70 124

Regenerable 8 14 22

Total FGD 236 432 668

Table 3-5. Average Size (MWe) of FGD Technology Systems (1998)

Technology United States Abroad World Total

Wet 472 330 377

Dry 261 152 199

Regenerable 350 171 236
32

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. 

SO2 Removal Efficiency

An estimate of SO2 removal performance of 

FGD processes can be obtained by examining 

the design SO2 removal efficiencies of these 

processes reported in the CoalPower3 

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 SO2 removal, up to 98 percent. 

However, most wet limestone systems appear 

to be designed for 90 percent SO2 removal. 

All LSFO systems installed after 1990 have 

design SO2 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 CoalPower3 

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 SO2 control.

Even though the median design efficiency for 

all units with wet limestone processes in 

CoalPower3 Database is 90 percent, it should 

be emphasized that advanced, state-of-the-art 

wet scrubbers are capable of routinely 

achieving SO2 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 CoalPower3 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 

CoalPower3 Database were used to determine 

median design SO2 removal efficiency. 
33

Table 4-1. Design SO2 Removal Efficiencies

FGD Technology Range of Design Efficiency, 

percent

Median Design Efficiency,a

percent

Wet Limestone Processes 52-98 90

LSD Processes 70-96 90

a 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 SO2 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 SO2 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
34

Figure 4-1. Design SO2 removal efficiencies for wet limestone and spray drying processes.

50

60

70

100

90

80

Wet Limestone

Spray Drying

1970s 1980s 1990s

Median

0

34
35

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. 
36

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 costeffectiveness 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.

Once-through Wet FGD Technology

At present, several technical options exist for 

upgrading the performance of existing 

installations using wet limestone processes. 

These options include:37

• increasing the sorbent amount used per 

mole of SO2;

• 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 

SO2 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 
37

requirements, are driving numerous 

conversions of existing older wet FGD 

systems to more advanced ones. These 

conversions are aimed at achieving improved 

SO2 removal efficiencies and/or waste 

minimization. Limestone wet FGD systems 

can be converted to MEL systems to increase 

SO2 removal efficiency. For example, an 

inhibited oxidation limestone wet scrubber 

designed for 85 percent SO2 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, SO2 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 SO2 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 Approach

Design 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
38

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 SO2 removal, sorbent 

utilization, and/or a particular system's 

operation. The increased SO2 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 ultrafine 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 
39

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 SO3 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 SO2 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 [(NH4)2SO4] 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 [SO3] and hydrogen chloride [HCl]) 

in addition to SO2. While HCl emissions can 

be reduced concurrently with SO2 emissions

using currently commercial FGD technology, 

the removal of SO3 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 (NH4)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 
40

navigable water or good rail access, and a 

location with high (NH4)2SO4 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 

HCl 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 SO2 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 [(NH4)2SO3] and 

ammonium bisulfite (NH4HSO3) to 

(NH4)2SO4 takes place in the reaction tank via 

injection of compressed air. (NH4)2SO4 

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, SO2/SO3, 

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 SO3 mist and droplet emissions.

(NH4)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.
41

Advantages to clear solution operation 

include:

• Discreet (NH4)2SO4 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

To maximize the by-product value, the 

(NH4)2SO4 material must be converted to 

larger granular crystals. To accomplish this, 

raw (NH4)2SO4 material is transferred to a 

compaction system. In this system, fresh feed 

(NH4)2SO4 material is mixed with the raw 

(NH4)2SO4 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. SO2 from the flue gas is absorbed in the 

spray tower by water according to the 

equation:

 (5 -1) 2 2 2 3SO + H O ↔ H SO

The H2SO3 is then reacted in a reaction tank 

with ammonia to form (NH4)2SO3 and

NH4HSO3:

2 ( ) (5 - 2) 2 3 3 4 2 3 H SO + NH ↔ NH SO

( ) 2 (5 - 3) 2 3 4 2 3 NH4HSO3 H SO + NH SO ↔

(NH4)2SO3 and NH4HSO3 are also oxidized 

in the absorber (forced oxidation) to form 

(NH4)2SO4 and NH4HSO4:

( ) ( ) (5 - 4) 2 2 4 2 4

1

4 2 3 NH SO + O ↔ NH SO

 (5 - 5) 2 2 4 4

1 NH4HSO3 + O ↔ NH HSO
42

The NH4HSO4 is neutralized in the presence 

of ammonia and water to form (NH4)2SO4:

( ) (5 - 6) 4 2 4 2

4 4 3 2

NH SO H O

NH HSO NH H O

+

+ + ↔

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 11 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 

SO2 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.
43

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.
44

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 singlevariable 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 singlevariable 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 TAG's 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.
45

Table 6-1. Sensitivity Analysis of LSFO Annual Operating Cost (baseline cost of 10.31 mills/kWh)

Variable, units Baseline Variable’s 

High Value

Variable’s 

Low Value

Cost for 

High Value 

of Variable, 

mills/kWh

Cost for Low 

Value of 

Variable, 

mills/kWh

High Value 

Difference, d 

%

Low Value 

Difference, e 

%

Capacity, 

MWe

500 2000 100 6.57 22.62 -36.3 119.4

Heat Rate, 

Btu/kWh

10,500 11,000 8,000 12.25 9.13 18.8 -11.4

Coal Sulfur 

Content, %

3.43 4.0 1.5 10.60 9.71 2.8 -5.8

Coal Heating 

Valuea, Btu/lb

11,922 14,000 10,500 9.56 10.92 -7.3 5.9

Air Heater 

Outlet, °F

300 360 280 10.45 10.26 1.4 -0.5

SO2 Removal, 

%

95 98 90 10.36 10.22 0.5 -0.9

L/G 125 160 60 10.36 10.22 0.5 -0.9

Slurry 

Concentration, 

% solids

15 20 10 10.31 10.31 0 0

Capacity 

Factor, %

65 90 40 8.04 15.41 -22.0 49.5

DBAc

Addition

no N/Tbyes N/T 10.20 N/T -1.1

Disposal Mode stacking landfill wallboard 12.51 10.23 21.3 -0.1

Absorber 

Material

alloy N/T RLCS N/T 9.66 N/T -6.3

No. of 

Absorbers

1 2 N/T 10.41 N/T 1.0 N/T

Reheat yes N/T no N/T 9.94 N/T -3.6

aCoal Data: Form EIA-767, DOE 28

bN/T=not tested

cDibasic Acid

dDifference = (Cost for High Value of Variable – 10.31) / 10.31 • 100%

eDifference = (Cost for Low Value of Variable – 10.31) / 10.31 • 100%
46

Table 6-2. Representative Values for LSFO Variables with Minor Cost Impacts

Variable Units Value Comments

Coal Heating Valuea Btu/lb 11,900 Baseline 11,922

Limestone Composition % CaCO3 95.3

SO2 Control Efficiency % 95

L/G gal/1,000 ft3

125 70 with DBA

Ambient Pressure in. Hg 29.4

Air Heater Outlet Temperature °F 300

Moisture in the Flue Gas % 14.0

Max Fan Capacity cfm 1,600,000 Either 2, 4, or 8 fans

Chimney Inlet Gas Temperature °F 127

a

 Not a minor impact; value is set to 11,900 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 (BMF), 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 BMF, 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:

 (6 -1)

32

% 1000 64

2 MW HR

HHV

Wt S

FR e SO • •













•

•

=

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 CaCO3 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. 
47

Figure 6-1. Schematics of LSFO system's equipment areas.

Limestone Water

Crushing

Station

Slurry

Preparation

Tank

Flue Gas

In

Reaction

Tank

Dew atering

Disposal

Absorber

Slurry

Bleed

Chimney

REAGENT FEED SO2

 REMOVAL WASTE/BY-PRODUCT

FLUE GAS HANDLING SUPPORT

EQUIPMENT

Flue

Gas

Out

47
48

CUECost adjusts the reagent feed ratio to 

ensure the CaCO3 present is sufficient to 

remove all the chlorine in the coal as CaCl2, 

in addition to the specified SO2 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 SO2

removal.

The BMF cost was estimated based on the 

limestone feed rate. Limestone composition 

(purity) has been fixed in this model at 95.3 

percent CaCO3, 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 (FRL) was estimated as:

 (6 - 2)

0.953

0.95

64

100 1.05

2

= • • • FRL FRSO

These parameters allow BMF to be estimated 

as follows:

 (6 - 3)

7118470

1000

68164.7

1000

494.55

1000

2.1128

1000

0.0034

&

2

4 3

B H DBA

L L

L L

F

C C

FR FR

FR FR BM

+ +

+ • +





























− •





























+ •





























= − •

where CB&H is the cost of the ball mill and 

hydroclones as given by:

1854902 (6 - 4)

2000

22412

2000

32.9

2

&

 +











•

 +











= •

L

L

B H

FR

FR C

and CDBA is the cost of the DBA tank as given 

by:

( )

 (6 - 5)

60

8.34 1 0.5

2000

0.95 20

364627

0.283

2

























• +

•

•

= •

SO

DBA

FR

C

CDBA was added only for LSFO systems with 

the DBA addition.

SO2 Removal Area

BMR cost (including absorbers, tanks, and 

pumps) – a third order polynomial on SO2

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 
49

was fixed at 70 for LSFO with DBA 

and, otherwise, was fixed at 125. 

CUECost default is L/G of 125 for 95 

percent SO2 removal.

BMR 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 Ffactor (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/106 Btu for lignite) were 

expected to have negligible impact on the 

estimated scrubber cost. Flue gas flow into 

the absorber (ACFM) was calculated as 

follows:

( )

( )

( )

( )

 (6 - 6)

0.209

0.04 0.209 0.04

100 6

100

528

460 295

60

9780

10

1000

6

















−

−

• • • + •

−

•

+

= • •

P

P

P P

MW HR

ACFM

e

where P is %O2 in the stack (9 percent O2 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 H2O 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:

 (6 - 7)

MODULE A

a p

R R

N PUMP N

BM BARE BSORBER

• + •

= +

 

where ABSORBER is the absorber cost equal 

to:

 (6 - 8)

1000

ABSORBER1 173978

0.5575

Na

ACFM  •











= •

or to:
50

 (6 - 9)

1000

ABSORBER 2 230064

0.5638

Na

ACFM  •











= •

for the RLCS or alloy material of 

construction, respectively (Na is number of 

absorbers).

The cost of pumps, PUMPS, was expressed 

as:

910.85 (6 -10)

0.5954

p

p

GPM N

N

F

PUMPS •

















= •

 

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 SO2 Removal Area 

(BARE MODULER) was calculated as 

follows:

1905302 (6 -11)

1000

34809

1000

188.2

1000

BARE MODULE 0.8701

2 2

2

2

3

+

















+ •

















•

−

















= •

SO SO

SO

R

FR FR

FR

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. H2O vacuum. Temperature at the 

fan inlet was fixed at 295 °F (CUECost 

method).

If the design incorporated reheat, BMG 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, BMG

cost was estimated as follows:

= BARE MODULE + ID FANS (6 -12) BMG G
51

The auxiliary cost of the Flue Gas Handling 

Area (BARE MODULEG) was calculated as:

420141 (6 -13)

1000

1

1266.4

1000

1

559693 0.2009

1000

1288.82

1000

0.1651

1000

238203 0.000012

1000

777.76

1000

BARE MODULE 0.1195

2

3 2

2

+

















•

+ •

















•

 + − •











•

 +











 − •











•

 + +











+ •













= − •

a

a

G

N

ACFM

N

ACFM ACFM

ACFM ACFM

ACFM

ACFM

where ACFM1 is flue gas flow rate out of the 

absorber.

The cost of fans (ID FANS) was calculated 

as:

ID FANS 91.24 (6 -14)

0.6842

f

f

N

N

ACFM

•

















= •

 

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 

SO2 mass flow rate for gypsum stacking was 

used based on CUECost. Moreover, a third

order polynomial on SO2 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 SO2 flow rate, SO2 removal 

efficiency (fixed at 95 percent), and CaCO3 in 

the limestone. All SO2 removed was 

assumed to be oxidized to form calcium 

sulfate dihydrate (gypsum). The BMW cost 

was estimated as follows:

BMW = BARE MODULEW + THICKENER (6 -15)

• For the Waste/By-product Handling 

System with gypsum stacking (BARE 

MODULEW1):

80700 (6 -16)

1000

1788

1000

4.0567

2

2

2

1

+

















•

+

















= − •

SO

SO

W

FR

FR

BM

 

• For the Waste/By-product Handling 

System with landfill (BARE 

MODULEW2):

773243 (6 -17)

1000

29091

1000

168.77

1000

0.325

2

2 2

3 2

2

+

















+ •

















− •

















= •

SO

SO SO

W

FR

FR FR

BM

• For the Waste/By-product Handling 

System with wallboard gypsum 

production (BARE MODULEW3):

1.25 (6 -18) 3 2 BMW = BMW •

The cost of thickener (THICKENER) was 

estimated as:

114562 (6 -19)

64 2000

172

THICKENER 9018.7 0.95

2

+

•

•

= • • FRSO
52

Support Equipment Area

The BME 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 BME 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:

BM BARE MODULE CHIMNEY (6 - 20) E = E +

For a BME with reheat, the cost of chimney 

(CHIMNEY 1) was estimated as:

CHIMNEY1 40208 1 (6 - 21) 0.3339= • ACFM

For a BME without reheat, the cost of 

chimney (CHIMNEY 2) was estimated as:

CHIMNEY 2 23370 1 (6 - 22) 0.3908= • ACFM

The auxiliary cost for Support Equipment 

Area was estimated as:

1993.8 1177674 (6 - 23)

BARE MODULE 0.0003 1.0677

2

3

• + • +

= • −

e e

E e

MW MW

MW

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 

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. Preproduction 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:

 (6 - 24)

100

1

100

B

1

100

A

100 100

1

1 2 3













 • +











• +













= • + + +

C

A A

TPC BM

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 deescalation of cost. CUECost includes time 

requirements for various size FGD 

installations.
53

Table 6-3. TCR Calculation Method

Cost Component Symbol / Calculation

Capital Cost BM = BMF + BMR +BMG + BMW + BME

Facilities + Engineering & HOa + Process Contingencies A = A1 + A2 + A3

Project Contingency B

Fee C

Total Plant Cost (TPC) TPC = BM * (1 + A ) * (1 + B) * (1 + C)

Financial Factor FTCE + FAFDC = D

Total Plant Investment TPI= TPC *D

Pre-production Cost + Inventory Capital E

Total Capital Requirement TCR = TPI + E

a

 HO = Home Office

Applying the FTCE and FAFDC appropriate to 

the unit size results in Total Plant Investment 

(TPI):

( ) (6 - 25) TPI TPC FTCE + FAFDC = •

 

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:

 (6 - 26)

12

&

12

&

1.02

INVENTORY

CF

VariableO M

FixedO M

TCR TPI

+

•

+

= • +

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.
54

Table 6-4. Financial Factors for FGD Construction, Constant Dollars

Unit Capacity Years to complete AFDC Factor TCE Factor

MWe < 160 1 0.0000 1.0000

160 # MWe < 400 2 0.0380 1.0000

400 # MWe < 725 3 0.0779 1.0000

725 # MWe < 1300 4 0.1199 1.0000

1300 # MWe < 2000 5 0.1640 1.0000

MWe = 2000 6 0.2104 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): 

 (6 - 27)

100

30 40 52 41.69041 MW 0.322307

e

• • •

• •

=

− MWe

OL

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:

8760 15 (6 - 29)

2000 3

= • •CF •

FR C

L

CaCO

where CF is capacity factor.

Cost of dibasic acid (unit price of dibasic acid 

at $430/ton):

8760 430 (6 - 30)

2000

20

2000

0.95

2

C = FR • • • •CF • DBA SO
55

Cost of disposal if the gypsum stacking 

method is selected ($6/ton):

 (6 - 31)

64 2000

172 6 8760 0.95

2

•

= • • • • • CDS CF FRSO

Cost of disposal if landfilling is selected 

($30/ton):

 (6 - 32)

64 2000

172 30 8760 0.95

2

•

= • • • • • CDL CF FRSO

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:

8760 2 (6 - 33)

64 2000

172 0.95

2

• • •

•

• •

=

FR CF

CREDIT

SO

Cost of steam (price of steam estimated at 

$3.50/1000 lb):

8760 3.5 (6 - 34)

855.14 1000

TER

• • •

•

STEAM = CF

Cost of electrical energy (power consumption 

for LSFO estimated at 2.0 percent) was 

estimated using the default CUECost power 

price of 25 mills/kWh:

( )

8760 25 (6 - 35)

1000

1000 0.8231 0.02

• • •

• •

= •

CF

MW POWER e

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.
56

Table 6-5. Model Validation Summary for LSFO FGD (1994 Dollars)

Plant Absorber Material/ 

Disposal

Unit Capacity, 

MWe

Absorbers Coal 

Wt%S

Model, 

$/kW

Reported, 

$/kW

Deviation,a

percent

Petersburg Alloy/landfill 239 1 3.5 400 317 +26.2

Cumberland RLCS/stacking 1300 3 4.0 164 200 -18.0

Conemaugh RLCS/wallboard 1700 5 2.8 174 195 -10.8

Ghent Alloy/stacking 511 3 3.5 213 215 -0.1

Bailly RLCS/wallboard 600 1 4.5 189 180 +5.0

Milliken RLCS/wallboard 316 1 3.2 368 348 +5.7

Navajo Alloy/landfill 750 2 0.75 226 236 -4.2

aDeviation=(Model-Reported)/Reported•100%

Figure 6-2. Comparison of model predictions with cost data for LSFO.

0

50

100

150

200

250

300

350

400

450

0 500 1000 1500 2000

Unit Capacity, MWe

Data ($/kW)

Model ($/kW)
57

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

module44,64,65. 

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.
58

Figure 6-3. Comparison of LSFO cost model to IPM model predictions for 2 to 4 percent sulfur coal.

0

100

200

300

400

500

600

0 500 1000 1500 2000

Unit Capacity, MWe

TCR, $/kW

IPM, 2%S ($/kW)

IPM, 3% ($/kW)

IPM, 4% ($/kW)

Model, 2%S ($/kW)

Model, 3%S ($/kW)

58

Model, 4%S ($/kW)
59

Table 6-6. “State-of-the-art” LSFO Design Decisions

Parameter Units Value

Single Absorber Size MWe 900

Absorber Diameter ft 60

DBA Additiona----- Yes

L/G gal/1000 ft370

O2 in Stack % 8

Material of Construction ----- Average of RLCS and alloy

SO2 Removal % 95

Flue Gas Temperature from Absorber °F 300

Flue Gas Velocity into Absorber ft/s 14

Inventory for Limestone days 30

Limestone Purity (CaCO3) % 95.3

Waste Disposal ----- Average of wallboard or gypsum stacking

Power Requirement % 2

Flue Gas Reheata---- Average of Yes and No

aYes/No decision only; no addition rate considerations
60

Figure 6-4. TCR predictions for 2 to 4 percent sulfur coal by LSFO SUSCM.

0

100

200

300

400

500

0 200 400 600 800 1000 1200 1400 1600 1800 2000

Unit Capacity, MWe

TCR, $/kW

SUSCM Model, 2% S

SUSCM Model, 3% S

SUSCM Model, 4% S

60
61

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.
62

Figure 6-5. Fixed O&M predictions for 2 to 4 percent sulfur coal by LSFO SUSCM.

62

0.0

2.0

4.0

6.0

8.0

10.0

12.0

14.0

16.0

18.0

0 200 400 600 800 1000 1200 1400 1600 1800 2000

Unit Capacity, MWe

Fixed O&M, $/kW

SUSCM Model, 2% S

SUSCM Model, 3% S

SUSCM Model, 4% S

- yr
63

Figure 6-6. Variable O&M predictions for 2 to 4 percent sulfur coal by LSFO SUSCM.

1.0

1.2

1.4

1.6

1.8

2.0

0 200 400 600 800 1000 1200 1400 1600 1800 2000

Unit Capacity, MWe

Variable O&M, mills/kWh

SUSCM Model, 2% S

SUSCM Model, 3% S

SUSCM Model, 4% S

63
64

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 SO2

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 (lb/h) and a power law component based 

on fresh lime slurry feed rate (gpm). The 

lime addition rate was determined by the 

uncontrolled SO2 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
65

Table 6-7. Sensitivity Analysis of LSD Annual Operating Cost (baseline value of 10.02 mills/kWh)

Variable, units Baseline Variable’s 

High Value

Variable’s 

Low Value

Cost for 

High Value 

of Variable, 

mills/kWh

Cost for Low 

Value of 

Variable, 

mills/kWh

High Value 

Difference,a

%

Low Value 

Difference, b 

%

Capacity, MWe 500 2000 100 4.76 18.77 52.5 87.3

Heat Rate, 

Btu/kWh

10,500 11,000 8,000 10.29 8.58 2.7 -14.4

Coal Sulfur 

Content, %

1.5 2.00 1.00 11.16 8.86 11.4 -11.6

Coal Heating 

Value, Btu/lb

11,922 14,000 10,500 9.14 10.78 -8.8 7.6

Air Heater 

Outlet, °F

300 360 280 10.10 9.99 0.8 -0.3

SO2 Removal, 

%

90 95 85 10.16 9.88 1.4 -1.4

Adiabatic 

Saturation 

Temp, °F

127 145 110 10.04 10.00 0.2 -0.2

Approach to 

Saturation, °F

20 50 10 10.05 10.01 0.3 -0.1

Recycle Slurry 

Solids, %

35 50 10 10.01 10.10 -0.1 0.8

# of Absorbers 2 3 N/Tc10.57 N/T 5.5 N/T

Absorber 

Material

RLCS Alloy N/T 10.69 N/T 6.7 N/T

a

(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
66

Table 6-8. Representative Values for LSD Variables with Minor Cost Impacts

Variable Units Value Comments

Coal Heating Value Btu/lb 11,900 Baseline 11,922

Lime Purity % CaO 90.0

SO2 Control Efficiency % 90.0

Ambient Pressure in. Hg 29.4

Air Heater Outlet Temperature °F 300

Moisture in the Flue Gas % 6.0

14.0

Before control device

After control device

Approach to Saturation °F 20

Adiabatic Saturation Temperature °F 127

Recycle Slurry Solids % 35

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:

 (6 - 36)

32

% 1000 64

2

MW HR

HHV

Wt S

FRSO e• •













•

•

=

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 SO2 flow rate is known, the Reagent 

Feed Area cost (BMF) may be estimated as 

follows:

(72338 ) (6 - 37)

3764611

1000

170023

0.3195 GPM

FR BM L

F

+ •













= • +

where FRL is the reagent feed rate:

 (6 - 38)

0.9

1 0.9

64

56 1.75

64

56 1.75

2 2

−

• • •

FRL = FRSO • • + FRSO

and GPM is slurry flow rate:

 (6 - 39)

60

1 0.3

8.34

0.3

1 0.3

56

74

56

74











 +

•

−

• + • •

=

L L

GPM

FR FR

F

SO2 Removal Area

SO2 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 
67

in CUECost to treat a maximum of 

300 MWe; larger units require 

multiple equal size absorbers.

The SO2 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:

( )

( )

( )

( )

 (6 - 40)

0.209

0.04 0.209 0.04

100 6

100

528

460 295

60

9780

10

1000

6

















−

−

• • • + •

−

•

+

= • •

P

P

P P

MW HR

ACFM

e

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 SO2 Removal Area cost 

(BMR), as shown below.

BM BARE MODULE SPRAY DRYERS (6 - 41) R = R +

• For the SO2 Removal System with RLCS 

construction, the cost of spray dryers 

(SPRAY DRYERS1) was calculated as:

 (6 - 42)

791896

1000

9246

1000

3.57

1

2

a

a a

N

N

ACFM

N

ACFM

SPRAY DRYERS

•

















+

















•

+ •

















•

− •

=

• For the SO2 Removal System with alloy 

construction, the cost of spray dryers 

(SPRAY DRYERS2) was calculated as:

 (6 - 43)

1080990

1000

12538

1000

4.85

2

2

a

a a

N

N

ACFM

N

ACFM

SPRAY DRYERS

•

















+

















•

+ •

















•

− •

=

where Na is the number of absorbers.

Auxiliary cost (BARE MODULER) was 

calculated as:

677421 % (6 - 44)

693335 % 214198

581877809 % 3653117 %

0.0966

3 2

−

+ •

•

















+ • +

• − •

=

Wt S

N

Wt S

Wt S Wt S

BARE MODULE

a

R

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 particulate 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 
68

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 particulate 

control device was assumed to be at 23 in. 

H2O vacuum and 147 °F. Flue gas exiting the 

fans was assumed to be at 1 in. H2O 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 particulate 

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 (BMG) was 

estimated as follows:

BMG = BARE MODULEG + ID FANS (6 - 45)

The area's auxiliary cost (BARE MODULEG) 

was estimated by the following equation:

 (6 - 46)

1000

3

2695.9

1000

2

4840.4

1000

1

47680

1000

15338

1000

1

1326.2

1000

1721.8

BARE MODULE

0.5 0.5

0.5 0.5576

0.683 0.7131





























 + •











+ •





























 + •











+ •

•





























 + •











•

=

ACFM ACFM

ACFM ACFM

N

ACFM ACFM

a

G

where ACFM1, ACFM2, and ACFM3 are 

flue gas flow rates at the exit from the 

absorber, particulate control device, and ID 

fans, respectively. Na is the number of 

absorbers.

The cost of ID fans (ID FANS) was 

calculated as:

 (6 - 47) 2ID FANS 91.24

0.6842

f

f

N

N

ACFM

•

















= •

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:

% 1026479 (6 - 48)

2051841884 % 1443163 2

• +

= • −

Wt S

BMW Wt S

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 (BME) 

included the chimney without reheat. The 

chimney cost (CHIMNEY) was based on the 
69

flue gas flow rate and was estimated as 

follows: 

23370 3 (6 - 49) 0.3908 CHIMNEY = • ACFM

Support Equipment Area cost (BME) was 

calculated as:

1354716.2 (6 - 50)

1.211 2704.2

2

CHIMNEY

BME MWe MWe

+ +

= − • + •

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 preproduction 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): 
70

( )

 (6 - 51)

100

30 40 52 18 25 2 278 • • •

= − • •

e

e

ln MW OL . . MW

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) (6 - 52)

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:

8760 65 (6 - 53)

2000

= • •CF •

FR C

L

CaO

where CF is capacity factor.

The cost of disposal ($30/ton) is: 

1000 0.1 (6 - 54)

64

129

30

2000

8760













• • + • • •

= • •

HHV

HR FR MW

C CF

SO e

DL

s

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):

( )

8760 25 (6 - 55)

1000

1000 0 007 • • ••

= • CF MW POWER .

e

As an annual expense, the components of 

variable O&M cost were adjusted for the 

capacity factor of the unit.

Validation

The 1995 EIA-767 browser database28 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 Unit Capacity, 

MWe

Coal S, wt % Number of 

absorbers

Reported Cost, 

$/kW

Model Cost, 

$/kW

Deviation,a %

H.L. Spurlock/2 508 3.6 4 189 222 17.5

Wyodak/1 362 0.8 3 172 203 18.0

North Valmy/2 267 0.5 3 231 205 -11.3

aDeviation = (Model – Reported) / Reported • 100%
71

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.
72

Figure 6-7. Validation of LSD cost model.

0

50

100

150

200

250

0 100 200 300 400 500 600

Unit Capacity, MWe

TCR, $/kW

Data ($/kW)

Model ($/kW)

72
73

Table 6-10. “State-of-the-art” LSD Design Decisions

Parameter Units Value

Single Absorber Size MWe 275

O2 in Stack % 8

Material of Construction ----- RLCSa

SO2 Removal % 90

Stoichiometry 1.4 for 2% S Coal

Flue Gas Temperature °F 300

Lime Inventory days 30

Lime Purity % 94

Lime Cost $/ton 50

Waste Disposal Cost $/ton 12

aRLCS = Rubber-lined Carbon Steel
74

Figure 6-8. LSD TCR predictions by LSD SUSCM.

74

0

50

100

150

200

250

300

350

400

0 500 1000 1500 2000

Unit Capacity, MWe

TCR, $/kW

SUSCM Model, 1%S

SUSCM Model, 2%S
75

Figure 6-9. LSD fixed O&M predictions by LSD SUSCM.

75

0.00

2.00

4.00

6.00

8.00

10.00

12.00

0 500 1000 1500 2000

Unit Capacity, MWe

Fixed O&M, $/kW-Yr

SUSCM Model, 1%S

SUSCM Model, 2%S
76

Figure 6-10. LSD variable O&M predictions by LSD SUSCM.

0.00

0.50

1.00

1.50

2.00

2.50

0 500 1000 1500 2000

Unit Capacity, MWe

Variable O&M, mills/kWh

SUSCM Model, 1%S

SUSCM Model, 2%S

76
77

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 stateof-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 (BMF) 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 BMF 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 BMF

cost to be estimated as follows:

3764611 72338 (6 - 56)

1000

170023 0.3195

GPM

L

F

F

FR

BM

• + + •

=

where FRL is reagent feed rate (lb/hr) and 

FGPM is slurry flow rate (gpm).

SO2 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 
78

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:

BARE MODULE ABSORBER PUMPS (6 - 57)

BM

R

R

+ +

=

• For SO2 Removal Area with alloy 

absorber construction:

 (6 - 58)

1000

230064 0 9

2

0 5638

a

.

N

ACFM

.

ABSORBER

 •











• •

=

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: 

 (6 - 59)

1000

173978 0 9

1

0 5575

a

.

N

ACFM

.

ABSORBER

 •











• •

=
79

The cost of pumps (PUMPS) was calculated 

as follows:

910.85 (6 - 60)

0.5954

p

p

GPM N

N

F

PUMPS •

















= •

where FGPM is slurry flow rate in gpm and Np

is the number of pumps.

The area auxiliary cost was estimated as 

follows:

 (6 - 61)

1905302

1000

34809

1000

188 2

1000

0 8701

0 825

2

2 2

3 2

































+

















+ •

















− •

















•

•

=

SO

SO SO

R

FR

FR

.

FR

.

.

BARE MODULE

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.

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. H2O 

vacuum. Temperature at the fan inlet was 

fixed in the model at 295 °F.

By fixing these design criteria, the BMG cost 

was estimated as follows:

BM BARE MODULE ID FANS (6 - 62) G = G +

where area auxiliary cost (BARE 

MODULEG ) was:

559693 (6 - 63)

1000

1288.82

1000

0.1651

1000

420141 0.000012

1000

1

1266.4

1000

1

238203 0.2009

1000

777.76

1000

0.1195

2

3

2

2

+













 + •











− •













+ + •

















•

•

+

















•

+ − •













 + •











− •

=

ACFM ACFM

ACFM

N

ACFM

N

ACFM

ACFM ACFM

BARE MODULE

a

a

G

and cost of fans (FANS) was:

91 24 (6 - 64)

0 6842

f

.

f

N

N

ACFM FANS . •

















= •

where Nf is the number of fans.
80

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 SO2 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 particulate 

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:

BARE MODULE THICKENER D & P (6 - 65)

BM

W

W

+ +

=

where:

1 25 (6 - 66)

773243

1000

29091

1000

168 77

1000

0 325

2

2 2

3 2

.

FR

FR

.

FR

.

BARE MODULE

SO

SO SO

W

•

































+

















+ •

















− •

















•

=

and
81

114562 (6 - 67)

64 2000

172 9018.7 0.95

2

+

•

• • •

=

FRSO

THICKENER

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:

BM BARE MODULE CHIMNEY (6 - 68) E = +

where:

 (6 - 69)

1993 8 1177674

0 0003 1 0667 0 825

3 2

















+ • +

• − •

•

=

e

e e

. MW

. MW . MW .

BARE MODULE

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:

23370 1 (6 - 70) 0.3908 CHIMNEY = • ACFM

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): 

 (6 - 71)

100 30 40 52

41.69041 MW 0.322307

e

• • •

= •

− MWe OL
82

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:

8760 50 (6 - 73)

2000

= • •CF •

FR C

L

CaO

where CF is the capacity factor.

The cost of disposal if gypsum stacking 

method is selected ($6/ton) was:

 (6 - 74)

64 2000

129 6 8760 0 95

2

•

C = • •CF • FR • . • DS SO

The cost of disposal for landfill ($30/ton) 

was:

 (6 - 75)

64 2000

129 30 8760 0.95

2

•

= • • • • • CDL CF FRSO

The cost of energy (energy consumption for 

MEL estimated at 1.05 percent) was 

estimated using the default CUECost energy 

price of 25 mills/kWh):

( )

8760 25 (6 - 76)

1000

1000 0.8231 0.0105

• • •

• •

= •

CF

MW POWER e

As an annual expense, the components of the 

variable O&M cost were adjusted for the 

capacity factor of the unit(s).

State-of-the-art Model

At this point, it is helpful to specify a "stateof-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-theart” 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. 
83

Table 6-11. “State-of-the-art” MEL Design Decisions

Parameter Units Value

Single Absorber Size MWe 275

O2 in Stack % 8

Material of Construction ----- Average of RLSC and alloy

SO2 Removal % 98

L/G gal/1000 ft340

Inventory for Lime days 30

Lime Purity (CaO) % 94

Sorbent Cost $/ton 50

Waste Disposal ----- wallboard

Power Requirements % 1.05

MEL/LSFO Capital Cost Ratio ----- 0.80-0.85

ID Fans Cost $

2

/3 of LSFO ID Fans Cost
84

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 Capacity Rangea

MWe

Capital Cost,

$/kW

Fixed O&M,

$/kW-Yr

Variable O&M,

mills/kWh

LSFOb100 - 1000 542 – 195 18 – 7 1.80 – 1.78

LSDc100 - 1000 363 – 140 12 - 4 2.24 – 2.24

MELd100 - 1000 384 – 238 16 – 8 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.
85

Figure 6-11. MEL TCR predictions by MEL SUSCM.

85

0

50

100

150

200

250

300

350

400

450

0 200 400 600 800 1000 1200 1400 1600 1800 2000

Unit Capacity, MWe

TCR, $/kW

SUSCM Model, 2%S

SUSCM Model, 3%S

SUSCM Model, 4%S
86

Figure 6-12. MEL fixed O&M predictions by MEL SUSCM.

86

0.00

2.00

4.00

6.00

8.00

10.00

12.00

14.00

16.00

18.00

0 200 400 600 800 1000 1200 1400 1600 1800 2000

Unit Capacity, MWe

Fixed O&M, $/kW-Yr

SUSCM Model, 2%S

SUSCM Model, 3%S

SUSCM Model, 4%S
87

Figure 6-13. MEL variable O&M predictions by MEL SUSCM.

87

0.00

1.00

2.00

3.00

0 500 1000 1500 2000

Unit Capacity, MWe

Variable O&M, mills/kWh

SUSCM Model, 2%S

SUSCM Model, 3%S

SUSCM Model, 4%S
88

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 coalfired 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 

(Hg0) in the high temperature regions of 

combustion devices. Hg0 is transformed into 

oxidized mercury (Hg++) as the flue gas cools. 

Therefore, the species predominantly present 

in flue gas include species of elemental Hg0

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 coalfired 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 coalfired 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 SO2

emissions is the effect that decreased 

emissions of SO2 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 μm in 

aerodynamic diameter (PM2.5).36 These 

aerosols are formed in the atmosphere in the 

presence of SO2 and other gases. Therefore, 

an increased scrubber SO2 removal 

efficiency, leading to lower SO2 emissions, 

may decrease the amount of PM2.5. Source 

emissions characterization is required to 

understand the fate of aerosol precursors 

(such as SO2) 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 SO2, NOX, and 

VOCs emitted from stationary sources. 72 

SO2 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 SO2. In this latter 
89

case, modern, state-of-the-art SO2 scrubbers, 

designed primarily for high efficiency control 

of SO2, 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-coalfired 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 Hg0, 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, Hg0 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 
90

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 Hg0 at 

the outlet of wet FGD systems. 77 The results 

of triplicate measurements revealed from 7.1 

to 38.5 percent increases of Hg0 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, Hg0

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 Hg0 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 

HCl 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 

Hg0.

79 Commercial solutions of chloric acid 

and sodium chlorate were capable of 

transferring 10 percent of Hg0 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 fullscale, 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 
91

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.
REFERENCES

92

 

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 National Air Quality and Emissions Trends 

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 Program Overview, U.S. Environmental 

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 Compliance Report, U.S. Environmental 

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5

 Keeth, R.J., P.A. Ireland, and P.T. Radcliffe, 

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 Chan, P.K., and G.T. Rochelle, “Modeling 

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Electric Utility Engineer's FGD Manual, 

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8

 Smith, K., M. Babu, and E. Goetz, "Review 

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 Moser, R.E., "Prospects for Inhibited 

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19 Jozewicz, W., G.T. Rochelle, and D.E. 

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20 Cole, J.A., J.C. Kramlich, W. R. Seeker, 

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21 Borgwardt, R.H., “Sintering of Nascent 

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22 Helfritch, D.J., P.L. Murphy, and E.G. 

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24 Taylor, H.S., and W. Nischt, "San Juan 

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25 Miller, E.J., "SO2 and NOX Retrofit Control 

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26 Binns, D.R., and R.G. Aldrich, "Design of 

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29 Klingspor, J.S., and C. Brogren, "LS-2, 

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30 “The Effects of Title IV of the Clean Air 

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31 Pearson, T.E., and J.C Buschmann, “High 

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32 Watson, G.B., L.J. Chaney, and W.F. 

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Tower Inlet Design," EPRI-DOE-EPA 

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33 Blythe, G., B. Horton, and R. Rhudy, 

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34 Brogren, C., R. Hakansson, K. Benton, and 

P. Rader, "Performance Enhancement Plates 

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CA, U.S. Department of Energy, Pittsburgh, 

PA, and U.S. Environmental Protection 

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Division, Research Triangle Park, NC, 1999. 

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35 Rogers, K.J., M. Young, and M. Hassibi, 

"Advances in Fine Grinding & Mill System 

Application in the FGD Industry," EPRIDOE-EPA Combined Utility Air Pollution 

Control Symposium: The Mega 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.

36 Gaikwad, R.J., and W.L. Boward, 

"Economic Evaluation of Controls for PM2.5 -

Case Study," EPRI-DOE-EPA Combined 

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Multi-pollutant 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. 

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37 Froelich, D., J. Landwehr, and D. 

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38 Feeney, R., "Upgrade Scrubbers to 

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39 Inkenhaus, W., L. Loper, M. Babu, and K. 

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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.

40 Straight, R.S., and J.D. Colley, 

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Symposium: The Mega Symposium: SO2

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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.

41 Keeth, P., P. Ireland, and P. Radcliffe, 

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42 DePriest, W., and J.M. Mazurek, "Key 

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43 Fukasawa, K. "Low Cost, Retrofit FGD 

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44 Smolenski, J.V., J.L. Murphy, and I.S. 

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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. 

TR-113187-V1.

45 Gohara, W.F., T.W. Strock, and W.H. Hall, 

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EPRI-DOE-EPA Combined Utility Air 

Pollutant Control Symposium: The Mega 

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Continuous Emission Monitors, EPRI, Palo 

Alto, CA, U.S. Department of Energy, 

Pittsburgh, PA, and U.S. Environmental 

Protection Agency, Air Pollution Prevention 

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46 Klingspor, J.S., and G.E. Bresowar, 

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47 Ellison, W., "Today's FGD Systems Satisfy 

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48 Kingston, W.H., D.K. Anderson, and W.P. 

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U.S. Department of Energy, Pittsburgh, PA, 

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Air Pollution Prevention and Control 

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49 Dudek, S.A., J. Rogers, and W.F. Gohara, 

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Pollution Prevention and Control Division, 

Research Triangle Park, NC, 1999. 

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50 Brogren, C., R. Fdhila, and P. Rader, 

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CA, U.S. Department of Energy, Pittsburgh, 

PA, and U.S. Environmental Protection 

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51 Weiler, H., and W. Ellison, "Wet GypsumYielding FGD Experience Using Quicklime 

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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.

52 Milobowski, M.G., "WFGD System 

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Combined Utility Air Pollutant Control 

Symposium: The Mega Symposium: SO2

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 

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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 

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Continuous Emission Monitors, EPRI, Palo 

Alto, CA, U.S. Department of Energy, 

Pittsburgh, PA, and U.S. Environmental 

Protection Agency, Air Pollution Prevention 

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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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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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58 Sedman, C. B., "Controlling Emissions 

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59 Ellison, W., "Worldwide Progress in 

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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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60 Borio, D., P. Rader, and M. Walters, 

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Volume 1: SO2 Controls, EPRI, Palo Alto, 

CA, U.S. Department of Energy, Pittsburgh, 

PA, and U.S. Environmental Protection 

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Division, Research Triangle Park, NC, 1999. 

TR-113187-V1.

61 Walsh, M.A., "New Marsulex Technology 

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EPRI-DOE-EPA Combined Utility Air 

Pollution Control Symposium: The Mega 

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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67 Personal communication from Bjarne 

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Wojciech Jozewicz of ARCADIS, USA, July 

31, 2000.

68 Personal communication from Manny Babu 

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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 

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72 Walker, K.D., K. Fritsky, S.J. Miller, G.L. 

Schelkoph, and G.E. Durham, "The Future of 

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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 M.J. 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, 

"Flue Gas Hg Measurements from Coal-Fired 

Boilers Equipped with Wet Scrubbers," in 

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AWMA, St. Louis, MO, June 1999.

76 Redinger, K.E., A.P. Evans, R.T. Bailey, 

and P.S. Nolan, "Mercury Emissions Control 

 

in FGD Systems," 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.

77 Brown, T.D., D.N. Smith, R.A. Hargis, Jr., 

and W.J. O'Dowd, "Mercury Measurement 

and Its Control: What We Know, Have 

Learned, and Need to Further Investigate," in 

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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 Wet FGD 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, 

"Process for Combined Control of Mercury 

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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80 Helfritch, D.J., P.L. Feldman, and E.G. 

Waugh, “The Use of a Circulating Fluid Bed 

for Mercury Adsorption and Particle 

Agglomeration,” in Proceedings: the Air & 

Waste Management Association: 92nd Annual 

Meeting & Exhibition, St, Louis, MO, June 

1999.

81 Helfritch, D.J., 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,” EPRIDOE-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.