EPA/600/R

- 00/093 November 2000


# CONTROLLING SO 2 

# 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

Prepa red for:



## U.S. Environmental Protection Agency Office of Research and Development Washington, D.C. 20460ii 

## 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. Environment al 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 activ ities 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 ou r 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 quali ty 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 t echnologies; 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.

Th is 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. Ti mothy Oppelt, Director National Risk Management Research Laboratoryiv 

## Abstract

Sulfur dioxide (SO 2 ) scrubbers may be used by electricity generating units to meet the requirements of Phase II of the Acid Rain SO 2 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 pre sents 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 applicati ons.



## 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 Acknowl edgements

................................ ................................ ................................ .................... iv List of Figures

................................ ................................ ................................ .......................... vii i 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 Li mestone 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 Rege nerable 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 Applicati ons

................................ ................................ ................................ .. 20 Current Application

................................ ................................ ................................ ....... 24

Chapter 4 Performance

................................ ................................ ................................ ...... 32 Introduction

................................ ................................ ................................ ................... 32 SO 2 Removal Efficiency

................................ ................................ ................................ 32 Energy Requirements

................................ ................................ ................................ .... 33 Applicability

................................ ................................ ................................ ................. 35vi

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 SO 2 Removal Area

................................ ................................ ................. 48 Flue Gas Handling Area

................................ ................................ ......... 50 Waste/By

- product Handling Area

................................ ........................... 51 Support Equipment Area

................................ ................................ ........ 52 Total Capital Requirement

................................ ................................ ................. 52 Ope ration 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 SO 2 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 SO 2 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

................................ ................................ ............................ 88vii

Introduct ion

................................ ................................ ................................ ................... 88 Once

- through Wet FGD

................................ ................................ ................................ 89 Dry FGD

................................ ................................ ................................ ....................... 90

References

................................ ................................ ................................ .......................... 92viii 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 Uni ted 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 SO 2 removal effic iencies 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 c oal

................................ ................................ ................................ ..................... 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

................................ ........................ 87ix List of Tables

3

- 1. Coal

- fired Electrical Generation Capacity (MW e ) Equipped with FGD

Technology (1998)

................................ ................................ ............................... 28 3

- 2. Total Capacity (MW e ) Equipped with Wet FGD Technology (1998)

.................... 29 3

- 3. Total Capacity (MW e ) Equipped with Dr y FGD Technology (1998)

..................... 29 3

- 4. Number of Installed FGD Technology Systems (1998)

................................ ......... 31 3

- 5. Average Size (MW e ) of FGD Technology Systems (1998)

................................ ... 31 4

- 1. Design SO 2 Removal Efficiencies

................................ ................................ ......... 33 5

- 1. Advanced Options for New Wet FGD Scrubbers

................................ .................. 37 6

- 1. Sensitivity Ana lysis 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 V alidation 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 Impact s

.............................. 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

............................ 84x List of Symbols

Symbol Meaning Unit

A&S Adminis tration 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 fa ns cfm

ACFM3 Flue gas flow out of particulate control device cfm

BARE MODULE E Support equipment area auxiliary cost dollars

BARE MODULE G Flue gas handling area auxiliary cost dollars

BARE MODULE R SO 2 removal area auxiliary cost dollars

BARE MODULE W Wa ste handling area auxiliary cost dollars

BM Capital cost for FGD system dollars

BM E Capital cost component for support equipment area dollars

BM F Capital cost component for reagent feed area dollars

BM G Capital cost component for waste handling area do llars

BM R Capital cost component for waste handling area dollars

BM W Capital cost component for waste handling area dollars

C B&H Cost of ball mill and hydroclones dollars

C CaCO3 Cost of limestone dollars

C CaO Cost of lime dollars

C DBA Cost of DBA tan k dollars

C DL Cost of disposal with landfilling dollars

C DS Cost of disposal with gypsum stacking dollars

CF Capacity factor %

CHIMNEY Cost of chimney dollars

CREDIT By

- product credit Dollarsxi

Symbol Meaning Unit

F AFDC Allowance for funds during co nstruction factor

------- F d F

- factor scf/10 6 Btu

F GPM Slurry flow rate gal/min

Fixed O&M Fixed operation and maintenance cost dollars

FR L Reagent feed rate lb/hr

FR SO2 SO 2 feed rate to the FDG System lb/hr

F TCE 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 ft 3

ML&M Maintenance, labor, and materials cost dollars

N a Number of absorbers

------- N f Number of fans

------- N p Number of pump s

------- 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

Variable O&M Variable operation and maintenance cost dollars

Wt%S Coal sulfur content Wt%xii Lis t 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 P rocess
MEL Magnesium

- Enhanced Lime Process

MW e Unit Electrical Generating Capacity NAAQS National Ambient Air Quality Standards

O&M Operation and Maintenance

PM 2.5 Particulate Matter Less than 2.5 : m (Aerodynamic Diameter)

RLCS Rubber

- Lined Carbon St eel

SUSCM State

- of

- the

- art Utility Scrubber Cost Model

TCR Total Capital Requirement

TPC Total Plant Cost

TPI Total Plant Investment

WESP Wet Electrostatic Precipitatorxiii Conversion Table

–English Units to SI Units

To Obtain From Multiply by

m ft 0.3048

m 2 ft 2 9.29 · 10 - 2

m 3 ft 3 2.83 · 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/m 2 in. Hg 345.311



## CHAPTER 1

INTRODUCTION

Combustion of sulfur

- containing fuels, such as coal and oil, results in sulfur dioxide (SO 2 ) formation. SO 2 emissions are known to cause detrimental impacts on human health and the environment. The major health concerns associated with exposure to high concentrations of SO 2 include breathing difficulties, respiratory illness, and aggravation of existing cardiovascular disease. In addition to the health impacts, SO 2 leads to acid deposition in the environment. This deposition causes

ac idification of lakes and streams and damage to tree foliage and agricultural crops. Furthermore, acid deposition accelerates the decay of buildings and monuments. While airborne, SO 2 and its particulate matter derivatives contribute visibility degradatio n.

Electric power generating units account for the majority of SO 2 emissions in the U.S. In 1998, these units contributed 64 percent of the national SO 2 emissions. 1 To mitigate SO 2 emissions from electric power generating units, the Acid Rain SO 2 Reduct ion Program 2
was established under Title IV of the Clean Air Act Amendments of 1990 (CAAA). This
two

- phase program was designed to reduce SO 2 emissions from the power generating industry.

Phase I of the Acid Rain SO 2 Reduction Program began on January 1, 1995, and ended December 31, 1999. In 1997, 423 power generating units, affected under Phase I, emitted 5.4 million tons of SO 2 (1.7 million tons below the allowable 7.1 million tons of SO 2 ). 3 Thus, the SO 2 emissions in 1997 reflect an output of 23 pe rcent below the allowable amount.

Phase II of the Acid Rain SO 2 Reduction Program began on January 1, 2000. The nationwide cap for SO 2 will be 9.48 million tons from 2000 through 2009. In 2010, the cap will be reduced further to 8.95 million tons, a leve l 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 SO 2 scrubbing) technologies.

This report presents a review of current FGD technologies. Following the introduction, Chapter 2 presen ts 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 technolo gy 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
resp onsible for developing and implementing SO 2 control strategies at sources, (2) persons involved in developing SO 2 and other regulations, (3) State regulatory authorities implementing SO 2 control programs, and (4) interested public at large. Moreover, pers ons 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 SO 2 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 comme rcial 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 tec hnologies can “conventionally” be classified as once

through and regenerable , depending on how sorbent is treated after it has sorbed SO 2
. 4 In

once

- through technologies, the SO 2 is permanently bound by the sorbent, which must be disposed of as a waste or utilized as a by

- product (e.g., gypsum). In regenerable technologies, the SO 2 is released from the sorbent during the regeneration step and may be further processed to yield sulfuric acid, elemental sulfur, or liquid SO 2 . The regenerated sorbent is recyc led in the SO 2 scrubbing step. Both once

- through and regenerable technologies can be further classified as wet

or dry . In wet processes, wet slurry waste or by

- product is produced, and flue gas leaving the absorber is saturated with moisture. In dry pro cesses, dry waste material is produced and flue gas leaving the absorber is not saturated with moisture.

Depending on process configuration and local market conditions at the plant site, once

through wet FGD processes can produce slurry waste or salable b y

- 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 cond ucted 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 ar e 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 is4 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
reg enerable 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), depen ding 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 by

products. All wet FGD processes discussed here are once

- through (i.e., non

- regenerable). A generalized flow diagram of a baseline wet FGD system is shown in Figure 2

- 2. SO 2 - containing flue gas is contacted with limestone slu rry 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 dis solution 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 to wer with flue gas flowing upwards, while limestone slurry is sprayed downwards by an array of spray nozzles. In the absorber, SO 2 is removed by both sorption and reaction with the slurry. Reactions initiated in the absorber are completed in a reaction ta nk, which provides retention time for finely ground limestone particles to dissolve and to react with the dissolved SO 2
.FGD TECHNOLOGY

5 Figure 2

- 2. Baseline wet FGD system.

Li m es tone W a ter

Cru s hing S tation S lurry

P reparation Ta nk F

lue Gas In F

lue Gas Out

Reac t ion Ta nk De w atering Di s po sal Ab s orber

S lurry Bleed Chimney

P rocess W a t er56

The overall reactions in the absorber and in the reaction tank can be summarized by:

1)

- (2 2 2 2 1 3 2

2 1 3 2 CO

O H CaSO O

H CaCO SO

+ • → + +

and

2)

- (2 2 2

2 2 4 2

3 2 2 1 2 CO

O H CaSO O H CaCO O SO

+ • →

+ + +

The complex chemistry summarized by the above equations involves SO 2 - CO 2 - H 2 O equilibrium relationships in the absorber, limestone dissolution, and sulfite/sulfate crystallization (occurring mostly in the

rea ction 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 SO 2 concentration and the excess air in the flue gas, as well as on slurry pH, some systems may be operated in the natural oxidation mode. However, for most applications, it is beneficial to control oxidation.

The dissolution and crystallization reactions in the reaction tank are, to a large extent, controlled by the pH of the liquid, which is a function of limestone stoichiometry (number of mols of Ca added per mol of SO 2 removed). Both pH and limestone stoichiometry are preset parameters for the operation of an absorber. Normally, the required stoichiometry of a limestone wet FGD system vari es from 1.01 to 1.1 moles of CaCO 3 per mole of SO 2 (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, CaSO 4 is easier to dewater than CaSO 3
. 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 eliminato r. 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 we ll 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., bette r 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 SO 2 c oncentration, 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 absorbers7

have a perforated tray added for the improvement of SO 2 capture. In such cases, the optimum flue gas velocity is determined by the tray design. For this type of absorber, excessive flue gas velocity will cause an absorber to "flood," wh ereas 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 SO 2 control is a packed absorber. Packed ab sorber 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 coul d 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 ft 3 of flue gas at actual conditions leaving the absorber. The amount of surface system available for the reaction with SO 2 is determined by L/G. For a counterflow spray absorber operated at a given flue gas flow r ate, L/G approximates the surface area of droplets and is one of the main design variables available to obtain a desired SO 2 removal in the absorber. The amount of available alkalinity for the reaction with SO 2 increases with the increasing L/G. L/G also affects the oxidation rate of sulfite/bisulfite reaction products in the absorber by affecting the absorption rate of O 2 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 SO 2 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
op erational range, increasing the amount of limestone added increases the amount of SO 2 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 SO 2 (increasing sulfur content of fuel) will decrease SO 2 removal efficiency by a wet absorber. This decreased efficiency is observed becau se increasing SO 2 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 s lurry 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 im prove 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 th e 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 scal ing 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 ox idation). 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 o xidation 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 d ewatering equipment. Nearly complete (99 percent plus) oxidation is required for a commercial quality by

- product.

This level of oxidation can be accomplished in a modern wet FGD system. However, the salability of the wet FGD by

- product (FGD gypsum) is also a function of the demand for gypsum. Depending on site

- specific conditions, LSFO may produce a salable by

product in the form of commercial grade gypsum that could be used for wallboard manufacturing. When salable gypsum is not attainable, dry FGD waste is piled (gypsum stacking) or landfilled. Gypsum stacking is the procedure where a gypsum slurry is sent to the stacking area, allowed to have the solids to separate from the water, and then removing the water and leaving the solids as a pile.

The solids handling system for LSFO consists of primary and secondary dewatering, solids modification unit, and ultimate waste disposal, regardless whether a part or all of the by

- product will be sold as commercial quality gypsum. The objective of primary dew atering 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
wate r 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 solids9

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 dispos al 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
concentr ation 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 quali ty 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 i s 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 th e 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
aggre gate, 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 com mercial quality gypsum is made, pelletization is employed. The selection of the ultimate disposal method is highly site

specific and depends on, among other factors, land availability, hydrogeology, and topography. In general, three options exist for the ultimate disposal of waste FGD solids: landfills, ponds, and gypsum stacks.

In addition to technical issues, several market issues are involved in the decision of wallboard manufacturers to use FGD gypsum. These market issues are presented below. Norma lly, 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 St ates 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 gyps um 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 gypsum10

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 (LSI O). 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 dec reasing 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: SO 2
co ncentration, 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 SO 2 /O 2 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 (Na 2 S 2 O 3 ) 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 consider ations, 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 limeston e 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 lim estone slurry, but is more expensive. The magnesium

enhanced lime process (MEL) is a variation of the lime process in that it uses a special type of lime, magnesium

- enhanced lime (typically 5

–8 percent magnesium oxide) or dolomitic lime (typically 20 pe rcent 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 m ore alkaline. Therefore, MEL is able to achieve high SO 2 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 SO 2 remov al. 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 oxid ation 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 gyps um 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 SO 2 . The chemistry of the process is similar to the LSFO ch emistry 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 condenser s. 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 SO 2
. The absorption of SO 2 takes place in an absorber, where seawater and flue gas are brought into close contact in a counter

current flow. The scrubber effluent flows to the treatment plant where it is air

- sparged to oxidize absorbed SO 2 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 backg round level normally is not detectable within even a short distance from the point of discharge.

Since the utilization of seawater for SO 2 scrubbing introduces a discharge to the ocean, it is necessary to make an assessment based on local conditions. Typ ically, 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, SO 2 - containing flue gas contacts alkaline (mo st 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), furnace12

sorbent injection process (FSI), and circulating fluidized bed process (CFB)]. The LSD and the CFB require dedicated absorber vessels for sorbent to react with SO 2 , while in DSI and FSI new hardware requirements are limited to sorbent deli very 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 SO 2 removals below those attainable with wet once

- th rough FGD.

Lime Spray Drying LSD for the control of SO 2 emissions is used for sources that burn low

- to medium

- sulfur coal, with occasional applications for coals with higher sulfur content. Some issues that limit the use of spray dryers with high

- sulfur coals include the potential impact of chloride contained in the coal on the spray dryer performance, and the ability of the existing particulate control device to handle the increased loading and achieve the required efficiency.

The LSD is shown schemati cally 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 nozz les 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 dispers ed lime slurry and SO 2 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 adiabati c saturation (from 10 to 15 ° C for coal

- derived flue gas) is required to achieve high SO 2 removal, complete saturation impairs operation of a spray dryer because of wet solids adhering to vessel walls and within the particulate collector. Primary reaction s in the spray dryer are as follows:

3)

- (2 ) ( 2 2 1 2 2 1 3 2 2 O H O H CaSO SO OH Ca + • → +

4)

- (2 2 ) ( 2 4 2 3 2 O H CaSO O H SO OH Ca •

→ + +

5)

- (2 4 2 2 1 3 CaSO 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 SO 2 removal may occur in the filter cake on the surface of bags. Dry solids from the particulate control device's hopper and from the bottom of the spray dryer are disposed of.

The extent of alkali usage in a spray dryer is limited by its available residence time for a
gas

- solid reaction. Typical residence time in a spray dryer is 8 to 12 seconds. In order to increase sorbent utilization, part of the dry solids from the bottom of the spray dryer and the particulate collector's hopper are sent to the recycle solids slurry tank. The recirculated stream (shown with a broken line in Figur e 2

- 3) contains partially reacted alkali from previous passes through the system. The additional exposure of a sorbent to SO 2 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 Chimney1314

Mass transfe r during a spray drying process occurs in two discrete phases: moist and dry. 17 During the moist phase, SO 2 diffuses from the bulk gas to the moisture layer on the surface of lime particles and reacts with dissolved lime. The reaction product precipitates on the surface of the lime particle. During the dry phase, SO 2 diffuses through the products of the lime and SO 2 reaction and causes a gas

- solid reaction with the unreacted core of lime particle.

Studies indicated that a majority of SO 2 capture in the s pray dryer occurs during the moist phase. Any increase in the duration of the moist phase would therefore increase the amount of captured SO 2 . Deliquescent salt additives sometimes are added to the lime slurry to be atomized in a spray dryer to achieve t his effect. A similar effect is achieved when spray dryers are used on coals with elevated chloride content.

Duct Sorbent Injection DSI for SO 2 emission control is intended to enable the control directly in the flue gas duct between the air preheater an d the particulate control device. Since no dedicated absorber vessel is required, the amount of hardware needed to control SO 2 is minimized for DSI.

DSI utilizes the contacting of finely dispersed sorbent with the flue gas. Sorbent used in DSI is typi cally 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 an d 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 alkal ine 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

re gion 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 s urface systems, 21 according to the reaction below:

6)

- (2 ) ( 2 2 O

H CaO OH Ca +

→

SO 2 in the flue gas reacts with the nascent CaO as given below:

7)

- (2 4 2 2 1 2 CaSO O SO CaO →

+ +

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. Sorben t 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

Chimney1516

Figure 2

- 5. Schematic of FSI. 

### Control Device Particulate

Disposal Boiler Flue Gas Out

Chimney Humidifier

Reactivation Reactor Sorbent Injection1617

Circulati ng 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 o f 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 c ontrol 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 SO 2 mass transfer conditions from the gas to the solid phase. This is achieved a s 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 MW e . 23 This process uses hydrated lime rather than the less expensive and less reactive limestone commonly use d 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 wit

h 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 (activat ed carbon). These processes are characterized by their product, a concentrated stream of SO 2 . 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 SO 2 in a wet scrubber where pretreated flue gas is contacte d with sodium sulfite solution. The product of the reaction is sodium bisulfite liquor heavily loaded with SO 2 . The liquor is subsequently regenerated in evaporators that crystallize sodium sulfite. Concentrated SO 2 is suitable for sulfuric acid product ion.

Magnesium Oxide In the magnesium oxide process, SO 2 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 SO 2 absorption in a scrubber . The absorbed product is dried and calcined in a kiln to regenerate magnesium oxide. SO 2 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

Fresh Sorbent Circulating Fluidized Bed Reactor Precollector1819

Sodium Carbonate In this process, SO 2 is contacted with a spray of sodium carbonate solution. Products of the reaction are sodium sulfite and sodium sulfate, which are reduced to sodium sulfide. Following the reaction of sodium sulfide with carbon dioxide and water, sodium carbonate is regenerated and hydroge n sulfide is converted to sulfur. 26

Amine The amine process involves absorption of SO 2 with an aqueous amine absorbent. The amine is regenerated thermally to release a concentrated water

- saturated SO 2 stream. SO 2 may then be treated by conventional tech nologies to produce sulfuric acid.

Dry Regenerable FGD Activated Carbon The activated carbon process adsorbs SO 2 on a moving bed of granular activated carbon. Activated carbon is thermally regenerated to produce a concentrated SO 2 stream. SO 2 may then b e 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 techn ologies 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 MW e 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 MW e 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 observe d throughout the world, as shown in Figure 3

- 2. With approximately 30,000 MW e of installed FGD capacity in 1980, the capacity has been increasing at an approximate rate of 100,000 MW e per decade. Similar to the trend in the United States, no significant increase in regenerable FGD capacity has taken place worldwide since the early '80s. Also, the rest of the world has seen a smaller percent of dry FGD

controlled capacity than the United States.

Since the wet FGD technology has historically dominated bot h 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 b y 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

- t o

- 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 proc esses 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 INSTALLEDe REGENERABLE

DRY

WET2122

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 INSTALLEDe REGENERABLE
DRY
WET2223

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 INSTALLEDMWe Mag LIME
LIME
SODIUM CARBONATE
LIMESTONE2324

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 t hroughout 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 m id '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, ma gnesium 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 th roughout 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 p rocess (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 increa se 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 t he impact of the Clean Air Act Amendments of 1990).


### Current Application Table 3

- 1 shows statistics describing the installation of FGD systems at fossil

- fuel

fired electric power plants through 1998. FGD systems were installed to control SO 2 emissions fro m over 226,000 MW e 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, a lmost 100,000 MW e 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 pe rcent 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 SO 2 removal efficiency, coupled with cost effectiveness, has been responsible for the overwhelming popularity of wet FGD technologies, particularly wet

- limestone

- based FGD technologies. While the earlier wet FGD systems produced only waste by

- product sludge , recent systems produce salable by

product gypsum. This has likely increased the attractiveness of wet FGD technologies. Limited application of dry FGD technologies, compared to wet FGD technologies, is likely the result of their higher reagent cost and limited choices for by

- product disposal.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 INSTALLEDMWe CIRCULATING FLUID BED
FURNACE INJECTION
SPRAY DRYING
DUCT SORBENT INJECTION2526

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 INSTALLEDMWe



## MgO
SODIUM SULFITE
OTHER2627

Figure 3

- 6. Percent shares (capacity) of the three FGD t echnologies 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%2728

Table 3

- 1. Coal

- fired Electrical Generation Capacity (MW e ) 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

T otal 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, li mestone 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 th ey 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 particula r 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 dr ying 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 domina nce 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 SO 2 removal is required.

Further understanding of recent FGD technology selections made by the U.S. electr icity 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 perce nt were equipped with LSFO, MEL, and LSD, respectively.29

Table 3

- 2. Total Capacity (MW e ) 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 (MW e ) Equipped with Dry FGD Technology (1998)

Process United States Abroad World Total
Sp ray 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 24,73530

Figure 3

- 7. Comparison of limestone and non

- limestone wet FGD applications. 

## Limestone

Non-Limestone U.S. 82,092 MW e



## 68.9% 31.1% Abroad 114,800 MW e



## 93.2% 6.8% World 196,892 MW e



## 83.0% 17.0%3031

Table 3

- 4 shows a dditional 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 techno logies.

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 MW e 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 regener able 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 17 4 348 522
Dry 54 70 124 Regenerable 8 14 22

Total FGD 236 432 668

Table 3

- 5. Average Size (MW e ) of FGD Technology Systems (1998)

Technology United States Abroad World Total

Wet 472 330 377
Dry 261 152 199 Regenerable 350 171 23632



## CH APTER 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.



### SO 2 

### Removal Efficiency An estimate of SO 2 removal performance of FGD processes can be obtained by examining the design SO 2 removal efficiencies of these processes reported in the CoalPower3 Database. Table 4

- 1 show s design SO 2 removal efficiencies for wet limestone and LSD processes. These data reflect that wet limestone systems have been designed for high levels of SO 2 removal, up to 98 percent. However, most wet limestone systems appear to be designed for 90 per cent SO 2 removal. All LSFO systems installed after 1990 have design SO 2 removal greater than 90 percent. The units with low design efficiencies are generally associated with plants burning low sulfur fuels. 28 Also, the units with the design efficiency at the low end of the range given in Table 4

- 1 are reported by the 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 SO 2 control.

Even though the median design efficiency for all units with wet limestone processes in CoalP ower3 Database is 90 percent, it should be emphasized that advanced, state

- of

- the

- art wet scrubbers are capable of routinely achieving SO 2 removal efficiencies of over 95 percent. The high velocity LSFO process, with state

- of

- the

- art design options, is

re portedly capable of removing more than 99.6 percent of SO 2 under test conditions. 29

As seen in Table 4

- 1, the range and median of SO 2 reduction efficiency at LSD installations are 70

- 96 and 90 percent, respectively. Spray dryers often achieve greater th an 90 percent SO 2 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 SO 2 removal efficiency of between 90 and 95 percent.

The performance of w et limestone and LSD processes has improved significantly over the period of their application. To investigate this improvement, the median design SO 2 removal efficiency was determined for the pertinent populations of wet limestone and LSD installations f

or 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 SO 2 removal efficiency.33

Table 4

- 1. Design SO 2 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 sul fur 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 SO 2 removal efficiencies, as well as ranges of repo rted design SO 2 removal efficiencies, for the wet limestone and LSD are shown in Figure 4

- 1. A steady increase of the design SO 2 removal efficiency can be noted for wet limestone and spray drying processes. This improvement may be due, in part, to the in creasing need to better control SO 2 emissions. However, the trends do reflect that the SO 2 removal efficiency for the processes considered has improved with time.


### Energy Requirements As described previously, once

- through wet FGD technology (and specific ally, 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 velocitie s 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 implemen ted 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 s urvey 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 SO 2 removal. Therefore, it is important to utilize a spray that has been atomized with in 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 SO 2 concentration in a34

Figure 4

- 1. Design SO 2 removal efficiencies for wet limest one and spray drying processes. 

## 50 60 70 100

90

80

Wet Limestone
Spray Drying

1970s 1980s 1990s Median

03435

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 MW e 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, c ontrol, 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 allo wed 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 dem onstrated 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

w ithout 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 i s 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 un its 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 con centration (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 th e 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 MW e . 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 fl ue 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 FG D 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 invol ved 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 a t improving the performance and cost

effectiveness of established processes, while others have focused on developing new processes. The initial part of this chapter discusses once

- through wet FGD technology advances. It discusses both advances that can b e 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

- th rough 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 SO 2 ;

• increasing the react ivity 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 m ass 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 ef

ficiencies of 95 percent or

more.

When considering the feasibility of upgrade scenarios, interrelations between increased SO 2 removal efficiency and many physical and technical parameters require a thorough evaluation. For example, the addition of

more s orbent 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 disp osal capacity. 38

The economics of FGD processes, affected by technical advances and regulatory37

requirements, are driving numerous conversions of existing older wet FGD systems to more advanced ones. These conversions are aimed at achieving improved SO 2 re moval efficiencies and/or waste minimization. Limestone wet FGD systems can be converted to MEL systems to increase SO 2 removal efficiency. For example, an inhibited oxidation limestone wet scrubber designed for 85 percent SO 2 removal at an L/G of 70 (ga l/1,000 ft 3 ) and 10 ft/s velocity has been converted to MEL lime. 39 Following the conversion, SO 2 removal efficiency increased to 96.7 percent at an L/G of 23 (gal/1,000 ft 3 ).

In another example of a vintage wet FGD system upgrade, conversion of an inhi bited 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 a nd 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
pr oviding high SO 2 removal and/or improving the operational efficiency of wet scrubbers while at the same time, reducing
cost.

Table 5

- 1. Advanced Options for New Wet FGD Scrubbers

Option Approach
Design large capacity modules

increased flue gas v elocity 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 disposal38

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 MW e 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 Sta tes 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 compu tational 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 velo city 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 g as/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 elimina te 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 fibergla ss, 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 determine d 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 SO 2 removal, sorbent utilization, and/or a particular system's operation. The increased SO 2 removal efficiency in the presence of DBA is a result of its buffering action (limiting t he 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 li mestone, 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 wet39

stack (no gas reheat) and a wastewater evaporation system. The latter is accompli shed 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 oxidat ion, 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

- prod uct Mg(OH) 2 can be produced optionally for in

- plant use or sale. This Mg(OH) 2 can be used for boiler injection for SO 3 control, to minimize air preheater fouling, and/or PM 2.5 - related stack emission.

MEL can offer some advantages over LSFO. It can oper ate with high SO 2 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 bee n under development. This process, wet ammonia FGD, has the potential to improve waste management in conjunction with providing SO 2 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 [(NH 4 ) 2 SO 4 ] by

- product that can be used as fertilizer. This process also has the potential for beco ming a promising option for units burning high sulfur coal, as it is also capable of removing other acid gases (e.g., sulfur trioxide [SO 3 ] and hydrogen chloride [HCl]) in addition to SO 2 . While HCl emissions can be reduced concurrently with SO 2 emissions using currently commercial FGD technology, the removal of SO 3 and control of sulfuric acid (H 2 SO 4 ) aerosol is not as straightforward. Depending on the type of FGD technology, a considerable portion of H 2 SO 4 aerosol may exit the stack as a respirable fine particulate emission and may cause a visible plume. 58

Ammonia scrubbing of SO 2 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 gen eration 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 abun dant 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

a mmonia, SO 2 , and oxygen in an absorber installed in an ammonia scrubbing system produces (NH 4 ) 2 SO 4 fertilizer. Recently, prices of ammonia are reported to have decreased considerably, 60 making the use of ammonia as a reagent much more economically favorab le compared to a few years ago. The best opportunities to apply ammonia scrubbing technology will likely be found at power plants in a proximity to40

navigable water or good rail access, and a location with high (NH 4 ) 2 SO 4 prices.

The ammonia scrubbing proc ess, 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 a

bsorption stage. The flue gas then enters the counterflow spray tower where it is contacted by a solution of (NH 4 ) 2 SO 4 liquor.

The ammonia is stored in a pressurized or refrigerated vessel and pumped as a liquid to a vaporizer. The vaporizer typically u ses 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 SO 2 remova l 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 [(NH 4 ) 2 SO 3 ] and ammoniu m bisulfite (NH 4 HSO 3 ) to (NH 4 ) 2 SO 4 takes place in the reaction tank via injection of compressed air. (NH 4 ) 2 SO 4 solution (10

- 25 wt percent dissolved solids) is bled from the absorber. Fresh makeup water required by the process is added to the absorber r eaction 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, SO 2 /SO 3 , 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

e missions. 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 SO 3 mist and droplet emiss ions.

(NH 4 ) 2 SO 4 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 (NH 4 ) 2 SO 4 (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 (NH 4 ) 2 SO 4 crystals are formed in a device specifically designed for that purpose, so that the size can be carefully controlled to meet th e required product specifications

• The monolithic crystals are not subject to attrition or dusting during shipping or handling

• The (NH 4 ) 2 SO 4 solution can be filtered prior to crystallization, thus eliminating any concern with solid contaminants (e.g. fly as h) 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 dewater ing 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 p articles (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 (NH 4 ) 2 SO 4 material must be converted to larger granular crystals. To accomplish this, r aw (NH 4 ) 2 SO 4 material is transferred to a compaction system. In this system, fresh feed (NH 4 ) 2 SO 4 material is mixed with the raw (NH 4 ) 2 SO 4 in a pug mill mixer. Finally, the material from the mixer is compacted into hard flakes subsequently discharged int o 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 (NH 4 ) 2 SO 4 from boiler f

lue gas is very similar to the chemistry of wet limestone FGD. SO 2 from the flue gas is absorbed in the spray tower by water according to the equation:

1)

- (5 3 2 2 2 SO

H O H SO ↔ +

The H 2 SO 3 is then reacted in a reaction tank with ammonia to form (NH 4 ) 2 SO 3 and NH 4 HSO 3 :

2)

- (5 ) ( 2 3

2 4 3 3 2 SO

NH NH SO H ↔ +

3)

- (5 2 ) ( 3 4 3 2 4 3 2 HSO

NH SO NH SO H ↔ +

(NH 4 ) 2 SO 3 and NH 4 HSO 3 are also oxidized in the absorber (forced oxidation) to form (NH 4 ) 2 SO 4 and NH 4 HSO 4 :

4)

- (5 ) ( ) ( 4 2 4 2 2 1 3 2 4 SO

NH O SO NH ↔

+

5)

- (5 4 4 2 2 1 3 4 HSO

NH O HSO NH ↔

+42

The NH 4 HSO 4 is neutralized in the pr esence of ammonia and water to form (NH 4 ) 2 SO 4 :

6)

- (5 ) ( 2

4 2 4 2

3 4 4 O

H SO NH O

H NH HSO NH

+ ↔

+ +

Because of the relatively high value of (NH 4 ) 2 SO 4 fertilizer, the economics of ammonia scrubbing improve as the sulfur content of the fuel increases. Consequently, ammonia scrubbing o ffers 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 b urned in many boilers. A recent study 61 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 w ith 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 (NH 4 ) 2 SO 4 fertilizer. An evaluation of the price of (NH 4 ) 2 SO 4 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 SO 2 remov al and aerosol

- free operation has recently been reported 60 for a 130 MW e 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 p rovide 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 (A PPCD) 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 MW e 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 determin e 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 model s 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 MW e 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 o f 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 m ills/kWh over alloy construction. Other variables were fixed to CUECost default values for the baseline LSFO, including 95 percent removal of SO 2
. This baseline LSFO retrofit requires one absorber serving a maximum size of 700 MW e . Thus defined, baselin e LSFO has an annual operating cost of 10.31 mills/kWh.

Sensitivity analyses were performed using the CUECost outputs resulting from single

variable perturbations from the baseline. The results of these analyses are summarized in Table 6

- 1. Perturbations in the variables were selected to span the range of realistic values (e.g., unit size ranged between 100 and 2000 MW e ). The high and low values of variables were selected and the corresponding costs were then determined for each single

variable perturbat ion. 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 (co st 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 f

ixed 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 operat ion 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 c apital 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 Differe nce, e

%

Capacity, MW e 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 Value a , Btu/lb 11,922 14,000 10,500 9.56 10.92

- 7.3 5.9

Air He ater Outlet, ° F 300 360 280 10.45 10.26 1.4

- 0.5

SO 2 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

DBA c Addition n o N/T b yes 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

a Coal Data: Form EIA

- 767, DOE 28

b N/T=not tested

c Dibasic Acid

d Difference = (Cost for High Value of Variable

– 10.31) / 10.31 • 100%

e Difference = (Cost for Low Value of Variable

– 10.31) / 10.31 • 100%46

Table 6

- 2. Representative Values for LSFO Variables with Minor Cost Impact s

Variable Units Value Comments
Coal Heating Value a Btu/lb 11,900 Baseline 11,922 Limestone Composition % CaCO 3 95.3

SO 2 Control Efficiency % 95

L/G gal/1,000 ft 3 125 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: reagen t feed, SO 2 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
represen ted as: Reagent Feed (BM F ), SO 2 Removal (BM R ), Flue Gas Handling (BM G ), Waste Handling (BM W ), and Support Equipment (BM E ). The estimation methods used for the five major equipment areas are described below.

The BM F , BM R , and BM W cost estimates were explic itly determined by the SO 2 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. SO 2 flow rate to the FGD system (FR SO2 ) was estimated from the
amount of sulfur in the coal as well as the coal burn rate at full load:

1)

- (6 32 64 1000 %

2 HR MW HHV S Wt FR e SO • •      

• •

=

where Wt%S is coal sulfur content (wt%), MW e is LSFO size (MW e ), HR is plant heat rate (Btu/kWh), and HHV is coal heating value (Btu/lb).

Reagent Feed Area The BM F c ost (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 SO 2 feed rate to the absorber, reagent addition rate, SO 2 removal

re quirement, and limestone CaCO 3 content. In CUECost (and in this simplified model), all the sulfur in the coal was assumed to be delivered to the FGD system as SO 2
.47

Figure 6

- 1. Schematics of LSFO system's equipment areas. Li me s tone W ater

Crushing S tation S lurry

P reparation Tank F

lue Gas
In

R

eaction Tank De w atering Disposal

Ab s orber

S lurry Bleed Chimney

REAGENT FEED SO 2 REMOVAL WASTE/BY-PRODUCT FLUE GAS HANDLING SUPPORT EQUIPMENT

F lue Gas Out4748

CUECost adjusts the reage nt feed ratio to ensure the CaCO 3 present is sufficient to remove all the chlorine in the coal as CaCl 2 , in addition to the specified SO 2 removal. However, chlorine removal has been eliminated from this model based on the assumption that it has a negligib le 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 i s, in turn, proportional to the rate of SO 2 removal.

The BM F cost was estimated based on the limestone feed rate. Limestone composition (purity) has been fixed in this model at 95.3 percent CaCO 3 , which is the default composition used in CUECost. The li mestone addition rate has been fixed in this model at 1.05 times the reagent feed ratio.

Reagent feed rate (FR L ) was estimated as:

2)

- (6 953

. 0 95

. 0
64 100 05

. 1
2 • • • = SO L FR FR

These parameters allow BM F to be estimated as follows:

3)

- (6 7118470 1000 7

. 68164 1000 55

. 494 1000 1128

. 2 1000 0034

. 0

& 2 3 4

DBA H B L

L L

L F

C C FR

FR FR FR BM

+ + + • +         

    

•

−         

    

• +         

    

•

− =

where C B&H is the cost of the ball mill and hydroclones as given by: 4)

- (6 1854902 2000 22412

2000 9

. 32 2

&

+      

• +

     

• =

L L H B

FR FR C

and C DBA is the cost of the DBA tank as given by:



### ( ) 5)

- (6

60 5

. 0 1 34

. 8 2000

20 95

. 0

364627 0.283

2 

  

 

   

 

+ • • •

• = SO

DBA FR C

C DBA was added only for LSFO systems with the DBA addition.

SO 2 Removal Area BM R cost (including absorb ers, tanks, and pumps)

–a third order polynomial on SO 2 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 MW e ; 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. H 2 O less than the inlet pressure (typical absorber inlet pressure drop). L/G49

was fixed at 70 for LSFO with DBA and, otherwise, was fixed at 125. CUECost default is L/G of 125 for 95 percent SO 2 removal.

BM R cost required estimation of flue gas flow through the LSFO system. The absorber cost was estimated bas ed on inlet flue gas flow rate and type of construction materials. The spray pump cost was estimated based on flue gas flow rates exhausting the absorber.

The flue gas flow was calculated in CUECost using the coal analysis in addition to unit size, heat rate, excess air, and air inleakage. This approach was analogous to computing the F

factor ( F

d ) for each fuel. As it was not considered practical to calculate an F d , for each fuel, gas flow was estimated using the methodology employed in 40CFR75 Appendi x F. An F d of 9,780 scf/10 6 Btu was applied for all coals, as the differences in coal rank (e.g., 9,860 scf/10 6 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)

- (6 209

. 0 04

. 0 209

. 0 04

. 0 6

100 100
528 295 460
60 9780
10 1000

6

       

−

− • + • • • − • + • • =

P

P
P P HR MW ACFM

e

where P is %O 2 in the stack (9 percent O 2 in the stack was assumed).

The pressure at the absorber inlet was fixed at 10 in. H 2 O gauge, the CUECost default. Ambient pressure was fixed at the CUECost default of 29.4 in. H g. Temperature of the flue gas entering the absorber might have been varying significantly for different units but was expected to have minimal impact on cost, based on the sensitivity analysis. Absorber inlet temperature was fixed in the model at 295 °F , resulting from 300 °F air heater outlet temperature used as the default in CUECost. The moisture fraction was assumed to be 6.0 percent H 2 O at the absorber inlet.

The cost of the spray pumps for the absorbers was estimated based on the absorber outlet f lue gas flow rate and the number of pumps (N p ) required. The N p 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, depe ndent on whether the design incorporated DBA additive. The gas flow rate was determined at 127 °F and at 9 in. H 2 O gauge (CUECost default). Moisture content was estimated at 14 percent H 2 O (CUECost calculated). CUECost estimated air addition at 2 moles oxygen for each mole of sulfite to be oxidized (CUECost default). For a typical SO 2 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 BM R cost, depen ding on the absorber construction material used and on the presence of DBA addition in the system as follows:

7)

- (6 A MODULE

p a R R N PUMP N BSORBER BARE BM

• + • + =

where ABSORBER is the absorber cost equal
to:

8)

- (6 1000 173978 1 ABSORBER 5575

. 0

a N ACFM

•      

• =

or to:50 9)

- (6 1000 230064 2 ABSORBER 5638

. 0

a N ACFM

•      

• =

for the RLCS or a lloy material of construction, respectively (N a is number of absorbers).

The cost of pumps, PUMPS, was expressed

as:

10)

- (6 85

. 910 5954

. 0

p p GPM N N F PUMPS •        

• =

where GPM is slurry flow rate (gpm) and N p is the number of pumps. The slurry flow rate varied, depending on w hether dibasic acid additive was selected.

Auxiliary cost for the SO 2 Removal Area (BARE MODULE R ) was calculated as follows:

11)

- (6 1905302 1000 34809 1000 2

. 188 1000 8701

. 0 MODULE BARE

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
p olynomial 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 BM G 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. H 2 O ga uge. 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 H 2 O 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 fan s was based on conditional tests of the smallest number option (2, 4, or 8) resulting in an individual fan capacity of less than 1,600,000 cfm. Inlet pressure for sizing fans was fixed at 12 in. H 2 O vacuum. Temperature at the fan inlet was fixed at 295 ° F (CUECost method).

If the design incorporated reheat, BM G 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, BM G cost was estimated as follows:

12)

- (6 FANS ID MODULE BARE + = G G BM51

The auxiliary cost of the Flue Gas Handling Area (BARE MODULE G ) was calculated as:

13)

- (6 420141 1000 1 4

. 1266 1000 1

2009

. 0 559693 1000 82

. 1288 1000 1651

. 0 1000 000012

. 0 238203 1000 76

. 777 1000

1195

. 0 MODULE BARE

2 2 3 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:

14)

- (6 24

. 91 FANS ID 6842

. 0

f f N N ACFM

•        

• =

where N f is the number of fans.

Waste/By

- product Handling Area The BM W cost (dewatering, disposal/storage, and washing)

- a second order polynomial on SO 2 mass flow rate for gypsum stacking was used based on CUECost. Moreover, a third order polynomial on SO 2 mass flow rate was used for landfill disposal or wallboard gypsum production. The cost of thickener was estimated as a linear function of the waste solids removal rate. The waste amount was estimated from a mass balance.

The BM W 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 SO 2 flow rate, SO 2 removal efficiency (fixed at 95 percent), and CaCO 3 in the limestone. All SO 2 removed was assumed to b e oxidized to form calcium sulfate dihydrate (gypsum). The BM W cost was estimated as follows:

15)

- (6 THICKENER MODULE BARE + = W W BM

• For the Waste/By

- product Handling System with gypsum stacking (BARE MODULE W1 ):

16)

- (6 80700 1000 1788

1000 0567

. 4

2 2 2

1

+        

• +        

•

− =

SO SO W

FR FR BM

• For the Waste/By

- product Handli ng System with landfill (BARE MODULE W2 ):

17)

- (6 773243 1000 29091 1000 77

. 168 1000 325

. 0

2 2 2 2 3

2

+

       

• +        

•

−        

• =

SO SO

SO W

FR FR

FR BM

• For the Waste/By

- product Handling System with wallboard gypsum production (BARE MODULE W3 ):

18)

- (6 25

. 1 2 3 • = W W BM BM

The cost of thickener (THICKENER) was estimated as:

19)

- (6 114562 2000 64 172 95

. 0 7

. 9018 THICKENER
2

+

• • •

• = SO

FR52

Su pport Equipment Area The BM E 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 BM E 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 wa s selected in the model at 127 °F, while the pressure was selected at 4 in. H 2 O gauge. The BM E cost was estimated as follows:

20)

- (6 HIMNEY MODULE BARE C BM E E + =

For a BM E with reheat, the cost of chimney (CHIMNEY 1) was estimated as:

21)

- (6 1 40208 1 CHIMNEY 3339

. 0 ACFM • =

For a BM E without reheat, the cost of chimney (CHIMNEY 2) was estimated as:

22)

- (6 1 23370 2 CHIMNEY 3908

. 0 ACFM • =

The auxiliary cost for Support Equipment Area was estimated as:

23)

- (6 1177674 8

. 1993 0677

. 1 0003

. 0 MODULE BARE

2 3

+ • + • − • =

e e e

E
MW MW MW

Total Capital Requirement Once the BM cost had been determined, it
was p ossible 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). Th e 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, inventor y cost, and pre

- production cost. Pre

production cost incorporated one

- twelfth of the projected annual O&M expense and 2 percent of the TPI estimate. Detailed calculations are described below.

Following the EPRI TAG approach, 5 percent for general facili ties, 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 wa s added to arrive at the Total Plant Cost (TPC). Using these CUECost defaults and adding yielded TPC for the model:

24)

- (6 100 1 100 B 1 100 A
100 100 1 3

2 1

      + •       + •       + + + • =

C A

A BM TPC

TPC could then be adjusted for financial factors dependent on the time required to complete the project. Allowance f or Funds During Construction Factor (F AFDC ) and Total Cash Expended Factor (F TCE ) are used to adjust TPC. F AFDC accounts for interest during construction and F TCE allows for de

escalation of cost. CUECost includes time requirements for various size FGD i nstallations.53

Table 6

- 3. TCR Calculation Method
Cost Component Symbol / Calculation Capital Cost BM = BM F + BM R +BM G + BM W + BM E Facilities + Engineering & HO a + Process Contingencies A = A 1 + A 2 + A 3 Project Contingency B
Fee C Total Plant Cost (T PC) TPC = BM * (1 + A ) * (1 + B) * (1 + C)
Financial Factor F TCE + F AFDC = 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 F TCE and F AFDC appro priate to the unit size results in Total Plant Investment (TPI):


### ( ) 25)

- (6 AFDC TCE F F TPC TPI +

• =

In regulatory cost determinations, it is usually preferable to assume constant dollars; e.g. no inflation. Such analysis should yield a F TCE equivalent to 1 (no in flation), and an F AFDC dependent on cost of capital without inflation. Applying an F AFDC 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 developme nt.

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: 26)

- (6 12 & 12 & 02

. 1

INVENTORY CF M VariableO M

FixedO TPI TCR

+

• + +

• =

where inventory capital (INVENTORY) is the cost of reagent required to meet the bulk storage requirement. A 60

- day limestone inventory was incorporate d (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 vari ables 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. Howe ver, 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 Facto rs for FGD Construction, Constant Dollars

Unit Capacity Years to complete AFDC Factor TCE Factor

MW e < 160 1 0.0000 1.0000

160 # MW e < 400 2 0.0380 1.0000

400 # MW e < 725 3 0.0779 1.0000

725 # MW e < 1300 4 0.1199 1.0000

1300 # MW e < 2000 5 0.1640 1.0 000

MW e = 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 sup port 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 l aw on the unit’s capacity and estimating the number of workers needed in combination with an operating labor rate ($30/hr):

27)

- (6 100 52 40 30 MW 69041

. 41 322307

. 0

e •

• • • • =

− e

MW OL Maintenance labor and materials (ML&M) cost was determined as a percentage (3 percent) of BM cost. Admi nistration and support (A&S) labor was estimated as a fraction of maintenance labor and materials plus operating labor, as given by the equation:


### ( ) 28)

- (6 & 4

. 0 3

. 0 & OL

M ML S A +

• • =

Variable O&M cost components were estimated as a sum of limestone, DBA, disposal, st eam, and electrical energy costs. Cost of limestone (unit price of limestone at $15/ton) was:

29)

- (6 15 8760 2000 3 •

• • = CF FR C L CaCO

where CF is capacity factor.

Cost of dibasic acid (unit price of dibasic acid at $430/ton):

30)

- (6 430 8760 2000 20
2000 95

. 0

2 •

• • • • = CF FR C SO DBA55

Cost of disposal i f the gypsum stacking method is selected ($6/ton):

31)

- (6 2000 64 172 95

. 0 8760 6
2 • • • • • • = SO DS FR

CF C Cost of disposal if landfilling is selected ($30/ton):

32)

- (6 2000 64 172 95

. 0 8760 30
2 • • • • • • = SO DL FR

CF C

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:

33)

- (6 2 8760 2000 64 172 95

. 0
2 • • •
• • • =

CF FR CREDIT

SO

Cost of steam (price of steam estimated at $3.50/1000 lb):

34)

- (6 5

. 3 8760 1000 14

. 855 TER

•

• •
• = CF STEAM

Cost of electrical energy (power consumption for LSFO estimated at 2.0 percent) was estima ted using the default CUECost power price of 25 mills/kWh:



### ( )

35)

- (6 25 8760 1000 8231

. 0 1000 02

. 0

• • • •

• • =

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 retrof it 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 mode l 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 t he 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/

Disp osal Unit Capacity, MW e 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 All oy/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

a Deviation=(Model

- Reported)/Reported • 100%

Figure 6

- 2. Comparison of mo del 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 predictions 63 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 Fi gure 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 algor ithms 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 desig n 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 s erving flue gas from units up to 900 MW e 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

- MW e module 44, 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 ass umed 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 discontin uities in capital cost curves reflect the addition of an absorber as unit capacity changes from less than 900 MW e to greater than 900 MW e and from less than 1800 MW e to greater than 1800 MW e . 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, M W eTCR, $/kW IPM, 2%S ($/kW) IPM, 3% ($/kW) IPM, 4% ($/kW)

M odel, 2%S ($/kW)

M odel, 3%S ($/kW)

M odel, 4%S ($/kW)5859

Table 6

- 6. “State

- of

- the

- art” LSFO Design Decisions

Parameter Units Value

Single Absorber Size MW e 900

Absorber Diameter ft 60

DBA Addition a

----- Yes

L/G gal /1000 ft 3 70

O 2 in Stack % 8

Material of Construction

----- Average of RLCS and alloy

SO 2 Removal % 95

Flue Gas Temperature from Absorber ° F 300

Flue Gas Velocity into Absorber ft/s 14

Inventory for Limestone days 30

Limestone Purity (CaCO 3 ) % 95.3

Waste Disposal

----- Average of wallboard or gypsum stacking

Power Requirement % 2

Flue Gas Reheat a

---- Average of Yes and No

a Yes/No decision only; no addition rate considerations60

Figure 6

- 4. TCR predictions for 2 to 4 percent sulfur coal b y LSFO SUSCM. 0 100 200 300 400 500

0 200 400 600 800 1000 1200 1400 1600 1800 2000 

## Unit Capacity, M W eTCR, $/kW SUSCM Model, 2% S

SUSCM Model, 3% S

SUSCM Model, 4% S6061

For comparison, the published average cost of 24 Phase I units 66 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 LS FO SUSCM predicts a LSFO TCR of $211/kW for a 500 MW e 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 S tation 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 an d the unit capacity (MW e ). 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
p rediction 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 th e 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

co nstant 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, M W eFixed O&M, $/kW SUSCM Model, 2% S

SUSCM Model, 3% S

SUSCM Model, 4% S-yr63

Fig ure 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, M W eVariable O&M, mills/kWh

SUSCM Model, 2% S
SUSCM Model, 3% S

SUSCM Model, 4% S6364



### Lime Spray Drying

Sensitivity Analysis Sensitivity analyses were performed to determine variables that have relatively minor impacts on FGD cost. The objective of these

ana lyses 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

- MW e unit wi th 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 MW e ). Other variables were fixed at CUECost default values for the baseline LSD, including 90 percent SO 2
r emoval 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 ca se.

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 cos t (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 SO 2 emission rate and the coal sulfur content. Fresh lime slurry feed rate was calculated for the lime addition rate at 30 percent solids, 1.3 specific gravity, and 90 percent lime purity.

The Reagent Feed Area cost (BM F ) was estimated based on the lime fe ed 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, which65

Table 6

- 7. Sensitivity Analysis of LSD Annual Operating Cost (basel ine 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, MW e 500 2 000 100 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

SO 2 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/T c 10.57 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%

c N/T = not tested66

Table 6

- 8. Representative Values for LS D Variables with Minor Cost Impacts

Variable Units Value Comments
Coal Heating Value Btu/lb 11,900 Baseline 11,922 Lime Purity % CaO 90.0

SO 2 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 t he maximum feed rate to the FGD system. As described earlier, the heating value was fixed at 11,900 Btu/lb in this model. The SO 2 flow rate can be estimated based on the coal sulfur content, unit capacity [MW e ], and heat rate [Btu/kWh] as follows:

36)

- (6 32 64 1000 %

2 HR MW HHV S Wt FR e SO • •      

• •

=

where Wt%S is coal sulfur content (wt%) MW e is LSD size, HR is plant heat rate (Btu/kWh), and HHV is coal heating value (Btu/lb).

Once the SO 2 flow rate is known, the Reagent Feed Area cost (BM F ) may be estimated as follows:



## ( ) 37)

- (6 72338 3764611

1000 170023

3195

. 0 GPM FR BM L F

• +       + • = where FR L is the reagent feed rate:

38)

- (6 9

. 0 9

. 0 1
64 56 75

. 1 64

56 75

. 1
2 2

− • • • +

• • = SO SO L FR FR FR

and GPM is slurry flow rate:

39)

- (6

60 3

. 0 1 34

. 8 3

. 0 3

. 0 1
56 74
56 74



     + • − • • + • =

L L GPM

FR FR F

SO 2 Removal Area SO 2 Removal Area cost (including spray dryers, tanks, and pumps)

- third order polynomial based on coal sul fur 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 limited67

in CUECost to treat a maximum of 300 MW e ; larger units require multiple equal size absorbers.

The SO 2 Removal Area cost (BM R ) 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:


### ( ) ( )

( ) ( ) 40)

- (6 209

. 0 04

. 0 209

. 0 04

. 0 6

100 100
528 295 460
60 9780
10 1000

6

       

−

− • + • • • − • + • • =

P

P
P P HR MW ACFM

e

The pressure at the absorber inlet was fixed at 12 in. H 2 O 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 SO 2 Removal Area cost (BM R ), as shown below.

41)

- (6 RS SPRAY DRYE E BARE MODUL BM R R + =

• For the SO 2 Removal System with RLCS construction, the cost of spray dryers (SPRAY DRYE RS1) was calculated as:

42)

- (6 791896 1000 9246 1000 57

. 3 1

2

a

a a N

N ACFM
N ACFM RS SPRAY DRYE

•        

+        

• • +        

• •

− = • For the SO 2 Removal System with alloy construction, the cost of spray dryers (SPRAY DRYERS2) was calculated as:

43)

- (6 1080990 1000 12538 1000 85

. 4 2

2

a

a a N

N ACFM
N ACFM RS SPRAY DRYE

•        

+        

• • +        

• •

− =

where N a is the number of absorbers.

Auxiliary cost (BARE MODULE R ) wa s calculated as:

44)

- (6 % 677421 214198

% 693335 %

3653117 % 581877809

0966

. 0 2 3

− • + •        

+

• + •

− • =

S Wt N S Wt S

Wt S Wt E BARE MODUL

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 part iculate 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 th e maximum gas flow rate per fan (1,600,000 cfm).

The Flue Gas Handling Area cost (BM G ) 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 I D fans. The68

flue gas exiting the absorber was assumed to be at 17 in. H 2 O 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. H 2 O vacuum and 147 °F. Flue gas exit ing the fans was assumed to be at 1 in. H 2 O gauge positive pressure and at 152 °F. The CUECost model adjusts flue gas flow rates to account for water evaporation and acid gas removal. For flue gas flow estimating purposes, all flue gas flows after the ab sorber 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 n umber 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. H 2 O.

Based on the assumptions presented above,
t he Flue Gas Handling Area cost (BM G ) was estimated as follows:

45)

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

The area's auxiliary cost (BARE MODULE G ) was estimated by the following equation:

46)

- (6 1000 3 9

. 2695 1000 2 4

. 4840 1000 1

47680 1000 15338 1000 1 2

. 1326 1000 8

. 1721 MODULE BARE

5

. 0 5

. 0 5576

. 0 5

. 0 7131

. 0 683

. 0

             

• +      

• +              

• +      

• + •              

• +      

• =

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. N a is the number of absorbers.

The cost of ID fans (ID FANS) was calculated as:

47)

- (6 2 24

. 91 FANS ID 6842

. 0

f f N N ACFM

•        

• =

where N f 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 (BM W ) was estimated as a function of coal sulfur content. Waste included fly ash and was presumed to b e sent to a landfill. BM W was estimated as follows:

48)

- (6 1026479 % 1443163

% 2051841884 2

+ • −

• = S Wt S

Wt BM W

Support Equipment Area Support Equipment Area cost (including electrical, water, and air)

- second order polynomial based on the unit capacity (MW e ).

- Cost of chimney

- power law based on the flue gas flow rate (ACFM3) exiting the ID fans.

The Support Equipment Area cost (BM E ) included the chimney without reheat. The chimney cost (CHIMNEY) was based on the69

flue gas flow rate and was estimated as follows:

49)

- (6 3 23370 3908

. 0 ACFM CHIMNEY •

=

Support Equipment Area cost (BM E ) was calculated as:

50)

- (6 2

. 1354716 2

. 2704 211

. 1 2

CHIMNEY MW MW BM
e e E
+ + •

+ •

− =

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 pos sible 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 befor e for LSFO, the adjustment results in Total Plant Investment (TPI) as described before in equation (6

- 25).

Since it is usually preferable to assume constant dollars in regulatory applications, a constant dollar analysis was done as explained before in the LSFO section.

Current dollar factors were used for validation, assuming that the published cost for TCR was in current dollars. Constant dollar factors were used in the subsequent model development.

TCR was determined by adding pre

production cost and inventory capital to TPI. CUECost estimates pre

- production cost at 2 percent of TPI plus one

- twelfth of the projected annual O&M (fixed plus variable adjusted for capacity factor) cost. Similar to considerations for LSFO, a 60 day lime inventory was inco rporated 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 Ta

ble 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

• a dministration 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 capacit y and estimating number of workers needed in combination with an operating labor rate ($30/hr):70 

### ( ) 51)

- (6 100 52 40 30 278 2 25 18 • • • • •

− = e
e MW ln MW

. . OL

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:



### ( ) 52)

- (6 M & ML 4 0 3 0 OL

. . S & A +

• • =

Variable O&M cost components were estimated as a sum of lime, disposal, fresh water, and energy costs. The cos t of lime (unit price of lime at $65/ton) was:

53)

- (6 65 8760 2000 •

• • = CF FR C L CaO

where CF is capacity factor.

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

54)

- (6 1

. 0 1000 64 129 30 2000 8760

     

• • • + • • •

• =

HHV HR MW FR CF C

e SO DL

s The cost of energy (energy consumption for LSD estimated at 0.7 percent) was estimated using th e default CUECost energy price of 25 mills/kWh):


### ( ) 55)

- (6 25 8760 1000 1000 007 0 • • • • • = 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 database 28 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 typ e 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 a ssumed 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, MW e Coal S, wt % Number of absorbers Reported Cost,

$/kW Model Co st,

$/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

a Deviation = (Model

– Reported) / Reported • 100%71

Figure 6

- 7. These results reflect that the simplified LSD cost mode l 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 cos t. 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 spe cific 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 MW e . 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 yiel ds 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 (MW e ) . 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, MW eTCR, $/kW

Data ($/kW)

Model ($/kW)727 3

Table 6

- 10. “State

- of

- the

- art” LSD Design Decisions

Parameter Units Value

Single Absorber Size MW e 275

O 2 in Stack % 8

Material of Construction

----- RLCS a

SO 2 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

a RLCS = Rubber

- lined Carbon Steel74

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, MW eTCR, $/kW
SUSCM Model, 1%S
SUSCM Model, 2%S75

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, MW eFixed O&M, $/kW-Yr

SUSCM Model, 1%S
SUSCM Model, 2%S76

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, MW eVariable O&M, mills/kWh

SUSCM Model, 1%S SUSCM Model, 2%S7677



### Magnesium

- enhanced Lime

General App roach The approach taken was to estimate the Magnesium

- enhanced Lime (MEL) system cost, both capital and O&M, based on the estimation methods previously described for LSFO and LSD. As described earlier, for costing purposes, MEL can be considered to be a combination of LSFO and LSD. The MEL cost was based on a retrofit presenting a medium difficulty. The derived algorithm was then further simplified by making state

of

- the

- art design decisions to build a cost model. TCR was estimated in the same

manner a s 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, SO 2 Removal, Flue Gas Handling, Waste Handling, and Support Equipment). Each major equipment a rea 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, SO 2 Removal, and Waste Handling Area cost estimates were expl icitly determined by the SO 2 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 11,90 0 Btu/lb. SO 2 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 (BM F ) co st (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 BM F 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 BM F cost to be estimated as follows:

56)

- (6 72338 3764611 1000 170023 3195

. 0 GPM L F

F FR BM

• + + • =

where FR L is reagent feed rate (lb/hr) and F GPM is slurry flow rate (gpm).

SO 2 Removal Area The SO 2 Removal Area (BM R ) 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 SO 2 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

co nstruction materials. Maximum78

absorber size was limited to 275 MW e ; 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
app lied 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. H 2 O less than the inlet pressure. L/G was fi xed at 40 consistent with the open tower design and 95 percent SO 2 removal.

The BM R 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

e stimated 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. Pres sure at the absorber inlet was fixed at 7 in. H 2 O 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 t o 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 a ssumed 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 ( n p ) req uired 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 SO 2 removal in an open spray tower. The gas flow rate was determined at 127 °F and at 6 in. H 2 O gauge. Moisture content was estimated at 14 percent.

These approximations allowed estimation of the BM R cost depending on the material of construction for the absorber as follows:

57)

- (6 PUMPS ABSORBER E BARE MODUL BM

R R
+

+ =

• For SO 2 Removal Area with alloy absorber construction:

58)

- (6 1000 9 0 230064 2

5638 0

a . N ACFM

. ABSORBER

•

     

• • =

where ACFM is flue gas flow at the absorber inlet in cfm and N a is the number of absorbers.

• For SO 2 Removal Area with RLCS absorber construction:

59)

- (6 1000 9 0 173978 1

5575 0

a . N ACFM

. ABSORBER

•

     

• • =79

The cost of pumps (PUMPS) was calculated as follows:

60)

- (6 85

. 910 5954

. 0

p p GPM N N F PUMPS •        

• =

where F GPM is slurry flow rate in gpm and N p is the number of pumps.

The area auxiliary cost was estimated as follows:

61)

- (6

1905302 1000 34809 1000

2 188 1000 8701 0
825 0

2 2 2 2 3



    

 

     

 

+

       

• +        

•

−        

•

• =

SO SO

SO R

FR FR

. FR

.

. E BARE MODUL

Flue Gas Handling Area Th e Flue Gas Handling Area (BM G ) 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 BM G 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. H 2 O gauge. The temperature of the gas e xiting 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. H 2 O vacuum. Temperature at the fan inlet was fixed in the model at 295 °F.

By fixing these design criteria, the BM G cost was estimated as follows:

62)

- (6 ID FANS E BARE MODUL BM G G + =

where area auxiliary cost (BARE MODULE G ) was:

63)

- (6 559693 1000 82

. 1288 1000 1651

. 0 1000 000012

. 0 420141 1000 1 4

. 1266 1000 1 2009

. 0 238203 1000 76

. 777 1000 1195

. 0

2 3 2 2

+      

• +      

•

−      

• + +        

• • +        

• •

− +      

• +      

•

− =

ACFM

ACFM ACFM
N ACFM N

ACFM ACFM ACFM E BARE MODUL

a a G

and cost of fans (FANS) was:

64)

- (6 24 91 6842 0

f .

f N N ACFM

. FANS •        

• =

where N f is the number of fans.80

Waste/By

- product Handling Area T he Waste/By

- product Handling Area (BM W ) 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 SO 2 mass flow rate). If forced oxidation is employed, system cost would be equivalent to LSFO gypsum stacking or wallboard by

- product options, also estimated as polynomials based on SO 2 mass flow rate. The cost of equipment components in this model was calculated as follows:

- Cost of thickener

- estimated wit h 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 co mponents 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 Wast e Handling Area cost, including the thickener.

Waste/By

- product Handling Area cost (BM W ) 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 pres umed to be dominated by filter cost. This model assumed 20 percent higher cost based on the SO 2 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 an d do not affect settling rates. The amount of particulate waste was based on inlet SO 2 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 SO 2 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 w as estimated at 10 percent of the waste handling system cost, including the thickener.

BM W cost was estimated as follows:

65)

- (6 & P D THICKENER E BARE MODUL BM

W W
+

+ =

where:

66)

- (6 5 2 1

773243 1000 29091 1000

77 168 1000 325 0

2 2 2 2 3

. FR FR

. FR

. E BARE MODUL

SO SO SO W

•      

 

     

 

+

       

• +        

•

−        

• =

and81 67)

- (6 114562 2000 64 172 95

. 0 7

. 9018
2 +

• • • • =

SO FR THICKENER

Bin and pugmill cost (D&P) was 10 percent of Waste Handling Area.

Support Equipment Area Support Equipment Area (BM E ) 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 cos t (BM E ) was estimated as follows:

68)

- (6 CHIMNEY E BARE MODUL BM E + =

where:

69)

- (6 1177674 8 1993 0667

1 0003 0 825 0 2 3 

      

+

• + •

− • • =

e e

e MW

. MW

. MW

.

. E BARE MODUL

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 mo del at 127 °F, while the pressure was fixed at 4 in. H 2 O gauge:

70)

- (6 1 23370 3908

. 0 ACFM CHIMNEY •

=

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 previou sly 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 Tot al 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 Require ment (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 pre diction.

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 co mposed 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 nee ded in combination with an operating labor rate ($30/hr):

71)

- (6 52 40 30 100 MW 69041

. 41 322307

. 0

e •

• • • = − e MW OL82

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 l abor and materials and operating labor as given by the equation below:


### ( ) 72)

- (6 M & ML 4

. 0 3

. 0 & OL S A +

• • =

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:

73)

- (6 50 8760 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:

74)

- (6 2000 64 129 95 0 8760 6
2 • • • • • • =

.

FR CF C SO DS

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

was:

75)

- (6 2000 64 129 95

. 0 8760 30
2 • • • • • • = SO DL FR

CF C

The cost of energy (energy consumpti on for MEL estimated at 1.05 percent) was estimated using the default CUECost energy price of 25 mills/kWh):



### ( )

76)

- (6 25 8760 1000 8231

. 0 1000 0105

. 0

• • • • • • =

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 "state

of

- the

- art" MEL system by which to measure the cost of possible future retrofits. Alternate design decisions may be made in the interest of reducing cost based on site specific conditions o r 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 MW e . The "state

- of

- the

- art" MEL scrubber used in this mo del incorporates RLCS or alloy absorber construction and salable gypsum. “State

- of

- the

- art” MEL design decisions are shown in Table 6

- 11.

Combining the equations developed earlier with these “state

- of

- the

- art” design decisions yields a model description of a “state

- of

- the

art” MEL FGD system.

MEL SUSCM TCR predictions for MEL are shown in Figure 6

- 11 for 2, 3, and 4 percent S coals. These predictions are based on units with a heat rate of 10,500 Btu/kWh and a capacity factor of 90 percent. MEL SUSCM
pre dictions 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 MW e 275

O 2 in Stack % 8

Material of Construction

----- Average of RLSC and allo y

SO 2 Removal % 98

L/G gal/1000 ft 3 40

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 Cost84

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 m ills/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 199 8 constant dollars for a 100 to 1000 MW e unit. As can be seen in Table 6

- 12, capital cost for LSFO used on a small unit (100 MW e ) is considerably higher than capital cost of MEL used on the same size unit. For a large unit (1000 MW e ), capital cost is

com parable 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 Range a

MW e Capital Cost,

$/kW Fixed O&M,

$/kW

- Yr Variable O&M,

mills/kWh

LSFO b 100

- 1000 542

–195 18

–7 1.80

–1.78

LSD c 100

- 1000 363

–140 12

- 4 2.24

–2.24

MEL d 100

- 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, SO 2 removal of 95 percent.

c 2.0 percent sulfur coal application, S O 2 removal of 90 percent.
d 4.0 percent sulfur coal application, SO 2 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, M W eTCR, $/kW

SUSCM Model, 2%S
SUSCM Model, 3%S
SUSCM Model, 4%S86

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, M W eFixed O&M, $/kW-Yr

SUSCM Model, 2%S
SUSCM Model, 3%S SUSCM Model, 4%S87

Figure 6

- 13. MEL variable O&M predictions b y MEL SUSCM.87



### 0.00 1.00 2.00 3.00

0 500 1000 1500 2000



## Unit Capacity, MW eVariable O&M, mills/kWh

SUSCM Model, 2%S

SUSCM Model, 3%S

SUSCM Model, 4%S88



## 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 SO 2 emissions. Mercury emissions from coal

fired power generation sour ces are reported to be almost 33 percent of the total anthropogenic emissions in the U.S. 69 In coal

- fired power generation, mercury is volatilized and converted to mercury vapor (Hg 0 ) in the high temperature regions of combustion devices. Hg 0 is transfor med into oxidized mercury (Hg ++ ) as the flue gas cools. Therefore, the species predominantly present in flue gas include species of elemental Hg 0 and Hg ++ . It follows that control of both of these mercury species is necessary to achieve total mercury emis sion control.

At present, the control of mercury emissions from coal

- fired boilers is not commercially practiced in the U.S. The combination of low mercury concentration and large flue gas volumes increases the difficulty and cost of controlling mercury emissions from coal

fired utility boilers compared to controlling mercury emissions from municipal waste combustors. 70 However, numerous studies have been conducted that reported some level of mercury emission control by the existing FGD processes. The ca pability of existing FGD processes to remove mercury from coal

fired flue gas is affected by the mercury species present. Because of mercury’s being the object of particularly strong concern due to its harmful effects on human health, the ability of the e xisting 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 SO 2 emissions is the effect that decreased emissions of SO 2 have on the formation of fine par ticulate aerosols. July 1997 revisions to the National Ambient Air Quality Standards (NAAQS) place emphasis on particulate matter less than 2.5 μ m in aerodynamic diameter (PM 2.5 ). 36 These aerosols are formed in the atmosphere in the presence of SO 2 and o ther gases. Therefore, an increased scrubber SO 2 removal efficiency, leading to lower SO 2 emissions, may decrease the amount of PM 2.5 . Source emissions characterization is required to understand the fate of aerosol precursors (such as SO 2 ) in the particl e formation process in the atmosphere. 71 While PM 2.5 can be produced directly by a variety of sources, it can also be produced by atmospheric reactions in the presence of SO 2 , NO X , and VOCs emitted from stationary sources. 72
SO 2 is a precursor for sulfur ic acid and sulfate secondary PM 2.5 particles. Sulfate accounts for approximately 47 percent of PM 2.5 in the eastern United States. 73 One strategy to control PM 2.5 emissions from stationary coal

- burning sources is to upgrade the existing particulate cont rol device. The alternative route may be to control PM 2.5 precursors, most notably SO 2 . In this latter89

case, modern, state

- of

- the

- art SO 2 scrubbers, designed primarily for high efficiency control of SO 2 , could provide an additional benefit by controlling PM 2.5 precursors.



### Once

- through Wet FGD A wide range of total mercury removal efficiency has been reported for once

- through wet FGD applications on bituminous

- coal

fired power generation units. Existing conventional wet scrubbers can remove
water

- solubl e Hg ++ compounds (e.g., mercuric chloride) from flue gas. However, a major part of Hg 0 , which is insoluble in water and the most volatile of the trace metal species, may pass through wet FGD and particulate matter control devices. 74 Therefore, should the control of mercury emissions be desired beyond the inherent control by once

- through wet FGD, Hg 0 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 me rcury 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 me rcury 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 stud y 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 correlatio n 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 ft 3 ). 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 ft 3 ) the mercury emissions for a system with tray scrubber were on the average 38 percent lower than these mea sured 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, with90

a scrubber configured as a tray tower and operated at an L/G of approximately 70 (gal/1,000 ft 3 ). 74

However, some sampling efforts have indicated an apparent re

- emission of Hg 0 at the outlet of wet FGD systems. 77 The results of triplicate measurements revealed from 7.1 to 38.5 percent increases of Hg 0 co ncentration across a wet FGD system operating on flue gas with inlet concentrations of from 2400 to 2900 ppm SO 2
.
As discussed before, should the control of mercury emissions by a wet FGD process alone be desired beyond the inherent control by the existin g once

- through wet FGD, Hg 0 would need to be converted by reagents or catalysts to a soluble form of mercury that could be collected.

Therefore, bench

- scale research and pilot studies are currently underway to more fully understand the oxidation of Hg 0 up stream 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 s tudy 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 Hg 0

. 79 Commercial solutions of chloric acid and sodium chlorate were capable of transferring 10 percent of Hg 0 into solution. Additionally, approximately 80 percent of the nitric oxide was removed. Further pilot

- scale evaluations continue to examine mercury speciation and to develop c ontrol options. 

### Dry FGD Similarly to wet FGD performance discussed above, a wide range of 55 to 96 percent reduction in mercury emissions has been shown with spray dryers installed on full

scale, bituminous

- coal

- fired boilers. 74 A significantly lower r eduction 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, comp ared 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 proce ss that is capable of additionally removing mercury is the CFB. Recently presented results of the pilot

- scale testing of a CFB process for mercury adsorption 80 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 zeolite 82

, is co

- injected along with the sorbent for SO 2 capture, then emissions of SO 2 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 is91

considered, at present, to be a niche application.

In summary, the amount of m ercury 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.REFERENCE S

92 1 National Air Quality and Emissions Trends Report, 1998, U.S. Environmental Protection Agency, Office of Air Quality Planning and Standards, Research Triangle Park, NC, March 2000; EPA

- 454/R

- 00

- 003 (NTIS PB2000

- 104960).

2 Program Overview, U.S. Envir onmental Protection Agency, Acid Rain Division, Washington, DC, 1997. Available at: http://www.epa.gov/acidrain/overview.html.

3 Compliance Report, U.S. Environmental Protection Agency, Acid Rain Division, Washington, DC, 1998. Available at: http://www. epa.gov/acidrain/cmprpt97/ cr1997.htm.

4 “Scrubber Myths & Realities,” White Paper prepared by Wet Gaseous Scrubber Division / Dry Gaseous Scrubber Division, Institute of Clean Air Companies, Inc., Washington, DC, May 1995.

5 Keeth, R.J., P.A. Ireland, a nd P.T. Radcliffe, “Economic Evaluations of 28 FGD Processess,” Palo Alto, CA, 1991. EPRI GS

7193.

6 Chan, P.K., and G.T. Rochelle, “Modeling of SO 2 Removal by Limestone Slurry Scrubbing: Effects of Chlorides,” in Proceedings: Eighth Symposium on Flue Ga s Desulfurization, Volume 1, EPA

- 600/9

- 84

- 017a (NTIS PB84

- 226638), Industrial Environmental Research Laboratory, Research Triangle Park, NC, July 1984.

7 Electric Utility Engineer's FGD Manual, Volume I

- FGD Process Design, U.S. Department of Energy, Off ice of Fossil Energy, Morgantown, WV, March 1996; DE

FG22

- 95PC94256

- 99. 8 Smith, K., M. Babu, and E. Goetz, "Review of 18 Months of Thioclear FGD Operation at AES Beaver Valley Cogeneration," EPRI

DOE

- EPA Combined Utility Air Pollution Control Symposium: The Mega Symposium: Volume 1: SO 2 Controls, EPRI, Palo Alto, CA, U.S. Department of Energy, Pittsburgh, PA, and U.S. Environmental Protection Agency, Research Triangle Park, NC, 1999. TR

- 113187

- V1.

9 Moser, R.E., "Prospects for Inhibited Oxidation FGD Sy stems," in Proceedings: 1993 SO 2 Control Symposium, Volume 3: EPA

- 600/R

- 95

- 015c (NTIS PB95

- 179248), Air and Energy Engineering Research Laboratory, Research Triangle Park, NC, February 1995.

10 Rochelle, G.T., D.R. Owens, J.C.S. Chang, and T.G. Brna, “Thio sulfate as an Oxidation Inhibitor in Flue Gas Desulfurization Processes: A Review of R&D Results,” in Ninth Symposium on Flue Gas Desulfurization, Volume 1, EPA

- 600/9

- 85

- 033a (NTIS PB86

- 138641), Air and Energy Engineering Research Laboratory, Research Tria ngle Park, NC, December 1985.

11 Polster, M., P.S. Nolan, and R.J. Batyko, "Babcock & Wilcox Technologies for Power Plant Stack Emissions Control," Paper No. BR

- 1571, U.S./Korea Electric Power Technologies Seminar Mission, Seoul, Korea, October 1994.

12 La ni, B.W., and M. Babu, "Phase II: The Age of High Velocity Scrubbing," EPRI

DOE

- EPA Combined Utility Air Pollutant Control Symposium: The Mega Symposium: SO 2 Control Technologies and Continuous Emission Monitors, EPRI, Palo Alto, CA, U.S. Department of Ene rgy, Pittsburgh, PA,REFERENCE S

93 and U.S. Environmental Protection Agency, Air Pollution Prevention and Control Division, Research Triangle Park, NC, 1997. TR

- 108683

- V2.

13 Benson, L., J. Garner, J.L. Murphy, M. Thompson, and C. Weilert, "Henderson Municipal Power & Light

-- A Low

- Cost Phase I Clean Air Act Retrofit,” EPRI

- DOE

EPA Combined Utility Air Pollutant Control Symposium: The Mega Symposium: SO 2 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.

14 Beeghly, J.H., M. Babu, and K.J. Smith, "Product Development of High Brightness FGD Gypsum and Magnesium Hydroxid e," EPRI

- DOE

- EPA Combined Utility Air Pollution Control Symposium: The Mega Symposium: Volume 1: SO 2 Controls, EPRI, Palo Alto, CA, U.S. Department of Energy, Pittsburgh, PA, and U.S. Environmental Protection Agency, Air Pollution Prevention and Control D ivision, Research Triangle Park, NC, 1999. TR

- 113187

- V1.

15 “Flue Gas Desulfurization by Seawater Scrubbing,” ABB Environmental Information No. 727, Staines, UK, 1997.

16 Williams, P., "Use of Seawater as Makeup Water for Water for Wet Flue Gas Desulfuriza tion Systems," EPRI

- DOE

- EPA Combined Utility Air Pollution Control Symposium: The Mega Symposium: Volume 1: SO 2 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.

17 Palazzolo, M.A., M.E. Kelly, and T.G. Brna, “Current Status of Dry SO 2 Control Systems,” in Proceedings: the Eighth Symposium on Flue Gas Desulfurization, Volume 2, EPA

- 600/9

- 84

- 017b (NT IS PB84

- 226646), July 1984.

18 Jozewicz, W., J.C.S. Chang, and G.T. Rochelle, “Reactivity of Ca(OH) 2 With SO 2 Under Duct Injection Conditions,” Presented at DOE Contractor’s Conference, Pittsburgh, PA, August 1990.

19 Jozewicz, W., G.T. Rochelle, and D.E. Stroud, “Novel Techniques for the Enhanced Utilization of Ca(OH) 2 Under Duct Injection Conditions,” Presented at DOE Contractor’s Conference, Pittsburgh, PA, July 1991.

20 Cole, J.A., J.C. Kramlich, W. R. Seeker, and M.P. Heap, “Activation and Reactivity o f Calcareous Sorbents Toward Sulfur Dioxide,” Env. Sci. Technol., 19, 1065 (1985).

21 Borgwardt, R.H., “Sintering of Nascent Calcium Oxide, Chem. Eng. Sci.

, 44:1 , 53 (1989).

22 Helfritch, D.J., P.L. Murphy, and E.G. Waugh, "The Pilot Scale Testing of a Cir culating 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. Dep artment of Energy, Pittsburgh, PA, and U.S. Environmental Protection Agency, AirREFERENCE S

94 Pollution Prevention and Control Division, Research Triangle Park, NC, 1999. TR

113187

- V3.

23 Graf, R.E., B.W. Huckriede, and G.G. Rochhausen, "10,000 Hours Commercial Operati ng Experience with Advanced

Design, Reflux Circulating Fluid Bed Scrubbing Employing Slaked Lime Reagent," 1995 SO 2 Control Symposium: Volume 2, EPRI, Miami, FL, March 1995.

24 Taylor, H.S., and W. Nischt, "San Juan Generating Station Limestone FGD Convers ion," EPRI

- DOE

- EPA Combined Utility Air Pollution Control Symposium: The Mega Symposium, SO 2 Control Technologies and Continuous Emission Monitors, EPRI, Washington, DC, August 1997.

25 Miller, E.J., "SO 2 and NO X Retrofit Control Technologies Handbook," EP RI, Palo Alto, CA, October 1985.

26 Binns, D.R., and R.G. Aldrich, "Design of the 100 MW Atomics International Aqueous Carbonate Process Regenerative FGD Demonstration Plant," in Proceedings: Symposium on Flue Gas Desulfurization: Volume II, EPA

- 600/7

- 78

- 0 58b (NTIS PB282091), Industrial Environmental Research Laboratory, Research Triangle Park, NC, March 1978.

27 Coal Power3, International Energy Agency, The Clean Coal Centre, London, UK (1998).

28 “Steam

- Electric Plant Operation and Design Report,” Form EI A

- 767, 1991 through 1995 Browser for Windows, Department of Energy, Washington, DC, 1997. 29 Klingspor, J.S., and C. Brogren, "LS

- 2, Two Years of Operating Experience," EPRI

DOE

- EPA Combined Utility Air Pollutant Control Symposium: The Mega Symposium: SO 2 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

- V 2.

30 “The Effects of Title IV of the Clean Air Act Amendments of 1990 on Electric Utilities: An Update,” DOE/EIA

- 0582 (97), U.S. Department of Energy, Energy Information Administration, Washington, DC (1997).

31 Pearson, T.E., and J.C Buschmann, “High Eff iciency Dry Flue Gas Desulfurization,” presented at: POWER

- GEN ’95, Anaheim, CA, December 1995.

32 Watson, G.B., L.J. Chaney, and W.F. Gohara, "Advanced, Low

- Pressure

- Drop, Tower Inlet Design," EPRI

- DOE

- EPA Combined Utility Air Pollution Control Symposium : The Mega Symposium: Volume 1: SO 2 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.

33 Blyth e, G., B. Horton, and R. Rhudy, "EPRI FGD Operating & Maintenance Cost Survey," EPRI

- DOE

- EPA Combined Utility Air Pollution Control Symposium: The Mega Symposium: Volume 1: SO 2 Controls, EPRI,REFERENCE S

95 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.

34 Brogren, C., R. Hakansson, K. Benton, and P. Rader, "Performance Enhancement Plates (PEP): Up to 20 percent reduction in power consumption of WFGD," EPRI

- DOE

- EPA Combined Utility Air Pollution Control Symposium: The Mega Symposium: Volume 1: SO 2 Controls, EPRI, Palo Alto, CA, U.S. Department of Energy, Pittsburgh, PA, and U.S. Environmental Protection Agency, Air Pollut ion Prevention and Control Division, Research Triangle Park, NC, 1999. TR

- 113187

- V1.

35 Rogers, K.J., M. Young, and M. Hassibi, "Advances in Fine Grinding & Mill System Application in the FGD Industry," EPRI

DOE

- EPA Combined Utility Air Pollution Control S ymposium: The Mega Symposium: Volume 1: SO 2 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 PM 2.5 - Case Study," EPRI

- DOE

- EPA Combined Utility Air Pollution Control Symposium: The Mega Symposium: Volume 2: NO X and Multi

- pollutant Controls, EPRI, Palo Alto, CA, U.S. Departmen t of Energy, Pittsburgh, PA, and U.S. Environmental Protection Agency, Air Pollution Prevention and Control Division, Research Triangle Park, NC, 1999. TR

- 113187

- V2. 37 Froelich, D., J. Landwehr, and D. Geschwind, "Compliance Options for Phase II of the Cl ean Air Act Amendments of 1990: A Look at Upgrading Existing FGD Systems," in Proceedings: the 1995 SO 2 Control Symposium, Miami, FL, March 1995.

38 Feeney, R., "Upgrade Scrubbers to Improve Performance." Power. Vol. 139, No. 8, pp. 32

- 37 (1995).

39 Inke nhaus, W., L. Loper, M. Babu, and K. Smith, "AEC Lowman Station FGD Conversion from Limestone to Magnesium

enhanced Lime Scrubbing," EPRI

- DOE

- EPA Combined Utility Air Pollutant Control Symposium: The Mega Symposium: SO 2 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.

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

, "Conversion of the 1600 MW Mill Creek Generating Station to Production of Commercial

- Grade Gypsum," EPRI

- DOE

EPA Combined Utility Air Pollutant Control Symposium: The Mega Symposium: SO 2 Control Technologies and Continuous Emission Monitors, EPRI, Palo A lto, 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, "Utility Response to P hase I and Phase II Acid Rain Legislation

-- An EconomicREFERENCE S

96 Analysis," in Proceedings: the 1995 SO 2 Control Symposium, Miami, FL, March 1995.

42 DePriest, W., and J.M. Mazurek, "Key Issues for Low

- Cost FGD Installation," in Proceedings: the Second International Conference on Energy and Environment: Transitions in East Central Europe, Prague, Czech Republic, October

- November 1994.

43 Fukasawa, K. "Low Cost, Retrofit FGD Systems," IEA Coal Research, Gemini House: London, UK, 1997.

44 Smolenski, J.V., J.L. Murphy, and I.S. Brodsky, "Tampa Electric Company Phase II FGD System Big Bend Station Units 1 & 2 Advanced Design

–Low Cost," EPRI

- DOE

EPA Combined Utility Air Pollution Control Symposium: The Mega Symposium: Volume 1: SO 2 Controls, EPRI, Palo Alto, CA, U.S. D epartment 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, "New Perspective of Wet Scrubber Flui d Mechanics in an Advanced Tower Design," EPRI

- DOE

- EPA Combined Utility Air Pollutant Control Symposium: The Mega Symposium: SO 2 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.

46 Klingspor, J.S., and G.E. Bresowar, "Advanced, Cost Effective Limestone Wet FGD," in Proceedings: the 1995 SO 2 Control Sym posium, Miami, FL, March 1995.

47 Ellison, W., "Today's FGD Systems Satisfy Retrofit Needs for 1990s," Power. Vol. 134, No. 2, pp. 101

- 106, 1991.

48 Kingston, W.H., D.K. Anderson, and W.P. Bauver II, "High Velocity Mist Elimination for Wet FGD Application ," EPRI

- DOE

- EPA Combined Utility Air Pollutant Control Symposium: The Mega Symposium: SO 2 Control Technologies and Continuous Emission Monitors, EPRI, Palo Alto, CA, U.S. Department of Energy, Pittsburgh, PA, and U.S. Environmental Protection Agency, Air P ollution Prevention and Control Division, Research Triangle Park, NC, 1997. TR

- 108683

- V2.

49 Dudek, S.A., J. Rogers, and W.F. Gohara, "CFD Models for Predicting Two

- Phase LFOR in Flue

- Gas Desulfurization Wet Scrubbers," EPRI

- DOE

- EPA Combined Utility Air Pollution Control Symposium: The Mega Symposium: Volume 1: SO 2 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.

50 Brogren, C., R. Fdhila, and P. Rader, "Computational Fluid Dynamic Modeling to Improve and Optimize the Design of WFGD of Spray Towers," EPRI

- DOE

- EPA Combined Utility Air Pollution Control Symposium: The Mega Symposium: Volume 1: SO 2 Controls, EPRI, Palo Alto, CA, U.S. Department of Energy, Pittsburgh, PA, and U.S. Environmental Protection Agency, Air Pollution Prevention and ControlREFERENCE S

97 Division, Research Triangle Park, NC, 1999. TR

- 113187

- V1.

51 Weiler, H., and W. Ellison, "Wet Gyps um

Yielding FGD Experience Using Quicklime Reagent," EPRI

- DOE

- EPA Combined Utility Air Pollutant Control Symposium: The Mega Symposium: SO 2 Control Technologies and Continuous Emission Monitors, EPRI, Palo Alto, CA, U.S. Department of Energy, Pittsburgh, P A, 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 Materials Cost Update," EPRI

- DOE

- EPA Combined Utility Air Pollutant Control Sympo sium: The Mega Symposium: SO 2 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.

53 Dille, E., K. Frizzell, and W. Shim, "Use of an Electrochemical Technique for Controlling Corrosion in FGD Reaction Tanks," EPRI

DOE

- EPA Combined Utility Air Pollution Control Symposium: The Mega Symposium: Volume 1: SO 2 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.

54 Chang, J.C.S., and T.G. Brna, "Enhanceme nt of Wet Limestone Flue Gas Desulfurization by Organic Acid/Salt Additives," in Proceedings: Tenth Symposium on Flue Gas Desulfurization, Volume 1, EPA

- 600/9

- 87

- 004a (NTIS PB87

- 166609), Air and Energy Engineering Research Laboratory, Research Triangle Par k, NC, February 1987.

55 Brogren, C., and J.S. Klingspor, "Impact of Limestone Grind on WFGD Performance," EPRI

- DOE

- EPA Combined Utility Air Pollutant Control Symposium: The Mega Symposium: SO 2 Control Technologies and Continuous Emission Monitors, EPRI, P alo 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.

56 Sarkus, T., P. Styf, and S. Vymazal, "Two Years of Outstan ding AGFD Performance, Pure Air on the Lake's Bailly Facility," Eleventh Annual International Pittsburgh Coal Conference, Pittsburgh, PA (September 1994).

57 Brown, G.N., K.E. Janssen, and P.A. Ireland, "Substantial Cost Reduction Realized for the First Am monia Scrubbing System," EPRI

- DOE

- EPA Combined Utility Air Pollutant Control Symposium: The Mega Symposium: SO 2 Control Technologies and Continuous Emission Monitors, EPRI, Palo Alto, CA, U.S. Department of Energy, Pittsburgh, PA, and U.S. Environmental Pr otection Agency, Air Pollution Prevention and Control Division, Research Triangle Park, NC, 1997. TR

- 108683

- V2.

58 Sedman, C. B., "Controlling Emissions from Fuel and Waste Combustion," Chemical Engineering, Vol. 106, No.1, pp. 82

- 88 (1999).REFERENCE S

98 59 Ellison , W., "Worldwide Progress in Ammonia FGD Application," EPRI

- DOE

EPA Combined Utility Air Pollution Control Symposium: The Mega Symposium: Volume 1: SO 2 Controls, EPRI, Palo Alto, CA, U.S. Department of Energy, Pittsburgh, PA, and U.S. Environmental Prote ction Agency, Air Pollution Prevention and Control Division, Research Triangle Park, NC, 1999. TR

- 113187

- V1.

60 Borio, D., P. Rader, and M. Walters, "Ammonia Scrubbing: Creating Value from SO 2 Compliance," EPRI

- DOE

- EPA Combined Utility Air Pollution Contr ol Symposium: The Mega Symposium: Volume 1: SO 2 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.

61 Walsh, M.A., "New Marsulex Technology Significantly Cuts Power Generation Costs," EPRI

- DOE

- EPA Combined Utility Air Pollution Control Symposium: The Mega Symposium: Volume 1: SO 2 Controls, EPRI, Palo Alto, CA, U.S. Department of Energy, Pittsburg h, 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. Kosmicki, D. Rhodes, and T. Waddell, "Coal Utility Environmental Cost (CU ECost) Workbook User's Manual, Version 1.0," EPA

- 600/R

- 99

- 056 (NTIS PB99

- 151938), June 1999.

63 “Analyzing Electric Power Generation under the CAAA,” Office of Air and Radiation, U.S. EPA, Washington, DC, April 1996, http://www.epa.gov/ardpublc/capihome/ca pi/ modeldoc/docnts.html

.

64 Schimmoller, B.K., “Balancing Compliance with Competition.” Power Engineering. Vol. 103, No. 10, pp. 22

- 28 (1999).

65 Jones, C.,"Meeting Compliance and Profit Goals Takes Ingenuity, Fortitude," Power, No. 7/8, pp. 63

- 65. 1999.

66 Ellerman, A.D., R. Schmalensee, P.L. Joskow, J.P. Montero, and E.M. Bailey, "Emissions Trading Under the U.S. Acid Rain Program, Evaluation of Compliance Costs and Allowance Market Performance," Center for Energy and Environmental Policy Research, MIT, C ambridge, MA, 1997.

67 Personal communication from Bjarne Rasmusseu of Niro A/S, Denmark, to Wojciech Jozewicz of ARCADIS, USA, July 31, 2000.

68 Personal communication from Manny Babu of Carmeuse Chimie Minerale, Pittsburgh, PA, to Ravi Srivastava, May 3, 2000.

69 Brenner, R., “Framework for the Future,” in Proceedings: the Conference on Air Quality, McLean, VA, December 1998.

70 Srivastava, R.K., C.B. Sedman, and J.D. Kilgroe, “Preliminary Performance and Cost Estimates of Mercury Emission Control Options for Electric Utility Boilers,” presented at 93 rd Annual Conference and Exhibition of the AWMA, Salt Lake City, UT, June 18

- 22, 2000.REFERENCE S

99 71 “Research Priorities for Airborne Particulate Matter II: Evaluating Research Progress and Updating the Portfolio,” Same t, J., Committee Chair; National Research Council, National Academy Press: 1999.

72 Walker, K.D., K. Fritsky, S.J. Miller, G.L. Schelkoph, and G.E. Durham, "The Future of Fine Particulate Matter and Hazardous Air Pollutants Control through the Utilization of GORE

- TEX(r) Membrane Filtration Technology," Presented at Air Quality: Mercury, Trace Elements, and Particulate Matter, McLean, VA: December 1998.

73 “Controlling Particulate Matter Under the Clean Air Act: A Menu of Options," State and Territorial Air Pollution Program Administrators (STAPPA), Association of Local Air Pollution Control Officials (ALAPCO); Washington, DC: July 1996.

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. Env ironmental 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 Proceedin gs: the Annual Meeting of the 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: SO 2 Con trol 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 Proceedings: the Air & Waste Management Association, 92nd Annual Meeting & Exhibition, St. Louis, MO, June 1999.

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 Cont rols, 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 Live ngood, 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, EPR I, Palo Alto, CA, U.S. Department of Energy, Pittsburgh, PA, and U.S. Environmental Protection Agency, Air Pollution PreventionREFERENCE S

100 and Control Division, Research Triangle Park, NC, 1999. TR

- 113187

- V3.

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: 92 nd 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 Mer cury Vapor Emissions Control,” EPRI

DOE

- EPA Combined Utility Air Pollution Control Symposium: The Mega Symposium, Volume 3: Particulate Controls, Air Toxics Control, and Poster Sessions, EPRI, Palo Alto, CA, U.S. Department of Energy, Pittsburgh, PA, and U .S. Environmental Protection Agency, Air Pollution Prevention and Control Division, Research Triangle Park, NC, August 1999.