WET FLUE GAS DESULFURIZATION TECHNOLOGY EVALUATION

PROJECT NUMBER 11311-001

PREPARED FOR

NATIONAL LIME ASSOCIATION

MAY 2006

PREPARED BY

55 East Monroe Street

Chicago, IL 60603-5780 USA
LEGAL NOTICE

This report was prepared by Sargent & Lundy LLC (Sargent & Lundy) expressly for National

Lime Association. Neither Sargent & Lundy nor any person acting on its behalf (a) makes any

warranty, express or implied, with respect to the use of any information or methods disclosed in

this report or (b) assumes any liability with respect to the use of any information or methods

disclosed in this report.
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CONTENTS

SECTION PAGE

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1. WET FLUE GAS DESULFURIZATION (FGD) DESCRIPTION .............................................. 1

1.1.1 LSFO Process.............................................................................................................................................................3

1.1.2 MEL Process ..............................................................................................................................................................5

1.2 Reagents and By-Products...........................................................................................................................................8

1.2.1 LSFO Process.............................................................................................................................................................8

1.2.2 MEL Process ..............................................................................................................................................................9

1.3 SO2 Removal Performance ........................................................................................................................................10

1.3.1 LSFO Process...........................................................................................................................................................10

1.3.2 MEL Process ............................................................................................................................................................10

1.4 Commercial Status .....................................................................................................................................................11

2. WET FGD TECHNOLOGY COMPARISON WITH DRY FGD TECHNOLOGY ................ 12

2.1 Process Advantages ....................................................................................................................................................12

2.2 Process Disadvantages................................................................................................................................................12

3. DESIGN BASIS ............................................................................................................................... 14

3.1 Specific Design Criteria..............................................................................................................................................14

3.2 System Components ...................................................................................................................................................16

3.2.1 Reagent Handling and Preparation Systems ............................................................................................................17

3.2.2 SO2 Removal System ...............................................................................................................................................17

3.2.3 Flue Gas System and Stack ......................................................................................................................................19

3.2.4 Gypsum Dewatering and Handling System .............................................................................................................20
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3.2.5 General Plant Support Systems ................................................................................................................................20

4. TECHNOLOGY APPLICATION CONSTRAINTS ................................................................... 23

4.1 Unit/Absorber Size .....................................................................................................................................................23

4.2 Coal Sulfur Content ...................................................................................................................................................23

4.3 Performance Expectations.........................................................................................................................................23

4.4 Reagent Utilization .....................................................................................................................................................23

4.5 By-Product Quality ....................................................................................................................................................24

4.6 Energy Consumption..................................................................................................................................................24

4.7 Retrofit Versus New Units .........................................................................................................................................24

4.8 Hg Removal.................................................................................................................................................................25

5. COSTS ANALYSES........................................................................................................................ 29

5.1 Capital Costs ...............................................................................................................................................................29

5.2 Operations and Maintenance Costs ..........................................................................................................................30

5.2.1 Fixed O&M Costs ....................................................................................................................................................30

5.2.2 Variable O&M Costs................................................................................................................................................31

5.3 Levelized Costs............................................................................................................................................................31
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EXHIBITS

NUMBER TITLE

5-1 Capital Cost for New Units Using Appalachian Coals – Wet FGD System

5-2 Capital Cost for Retrofit Units Using Appalachian Coals – Wet FGD System

5-3 Fixed and Variable O&M Cost for New Units – LSFO FGD System

5-4 Fixed and Variable O&M Cost for Retrofit Units – LSFO FGD System

5-5 Fixed and Variable O&M Cost for New Units – MEL FGD System

5-6 Fixed and Variable O&M Cost for Retrofit Units – MEL FGD System

5-7 Levelized Costs for New Units – LSFO FGD System

5-8 Levelized Costs for Retrofit Units – LSFO FGD System

5-9 Levelized Costs for New Units – MEL FGD System

5-10 Levelized Costs for Retrofit Units – MEL FGD System

APPENDIX

NUMBER TITLE

A Auxiliary Power Summary

B Detailed Description of MEL Process Chemistry
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REPORT PREPARED, REVIEWED, AND APPROVED BY SARGENT & LUNDY LLC:

Prepared by: May 2006

Raj Gaikwad Date

Technical Advisor

Reviewed by: May 2006

Steve Katzberger Date

Technical Advisor

Approved by: May 2006

William DePriest Date

Project Director
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1. WET FLUE GAS DESULFURIZATION (FGD) DESCRIPTION

Wet FGD technology, which is based on using limestone or lime as a reagent, is the scrubbing process that has

been the FGD technology most frequently selected for sulfur dioxide (SO2) reduction from coal-fired utility

boilers. The wet FGD flue gas treatment system is typically located after removal of particulate matter from the

flue gas either by a baghouse or by an electrostatic precipitator. The cleaned gas is discharged directly to the stack

with no further treatment such as reheat. Wet FGD processes are considered commercially mature technologies

and are offered by a number of suppliers. This report presents the results of Sargent & Lundy’s evaluation of both

limestone-based (limestone with forced oxidation -LSFO) and lime-based (magnesium-enhanced lime [MEL] with

forced oxidation) FGD technologies.

Flue gas is treated in an absorber by passing the flue gas stream through a limestone or lime slurry spray. In

typical absorber designs, the gas flows upward through the absorber countercurrent to the spray liquor flowing

downward through the absorber. However, other designs are also available, including co-current and

countercurrent designs, and designs where the gas is forced through the liquor in a froth-type bubbling absorber. 

In a typical design, slurry is pumped through banks of spray nozzles used to create fine droplets to facilitate

intimate and uniform contact with the flue gas. The droplets absorb SO2 from the gas, facilitating the reaction of

the SO2 with the reagent in the slurry. Hydrogen chloride and other soluble species (e.g. HF, Hg++, etc.) present

in the flue gas are also absorbed in the spray droplets, causing an accumulation of chloride ions (and other

absorbed species) in the process liquid. Some of the water in the spray droplets evaporates, adiabatically cooling

the gas from approximately 300°F at the inlet to 125°F-140°F, and saturating the flue gas with water. The

desulfurized flue gas passes through mist eliminators to remove entrained droplets before the flue gas is discharged

into the stack.

In most wet FGD systems, the specific SO2 collection efficiency required is achieved by selecting a combination

of design features for the system. For example, one feature is the quantity of liquid sprayed relative to flue gas

flow and is referred to as the liquid-to-gas (L/G) ratio. Higher L/G ratios improve SO2 removal by exposing the

gas to more absorbing liquor. However, higher L/G ratios also consume more power, and this design feature must

be factored against other important design features, including such design parameters as absorber velocity,
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absorber pressure drop and quantity/type of reagent. An important goal of this study will be to quantify the impact

that selection of the reagent (limestone or lime) will have on the design of the system and the resultant capital and

operations & maintenance (O&M) costs.

After contacting the gas, the slurry collects in the bottom of the absorber in a reaction tank. The slurry is agitated

to prevent settling. Limestone or lime consumed in the process is replenished by adding fresh limestone or lime

slurry to the reaction tank. 

In the LSFO process, the slurry is also aerated in the reaction tank to oxidize calcium sulfite hemihydrate (CaSO3 •

1⁄2 H2O) to calcium sulfate dihydrate (CaSO4 • 2H2O), or gypsum, which precipitates. This is where the term

"forced oxidation" originates and it distinguishes this process from older, more troublesome limestone-based

"natural oxidation" technology. The oxidized slurry is then recirculated to the spray headers. A portion of the

slurry is withdrawn to remove the precipitated gypsum. Typically, this slurry is dewatered in a two-stage process

involving a hydroclone and vacuum filter system to produce a gypsum cake for disposal or sale. Water removed

from the gypsum slurry is returned to the process. A portion of this water is removed from the system as

wastewater to limit accumulation of chloride salts and other undesirable constituents in the process liquid. 

In the MEL process, the slurry is aerated for the same reason, but in a separate tank, ultimately producing a

gypsum cake similar to the LSFO process. Water removed from the gypsum and soluble magnesium salts

(sulfites) are recycled to the process with a portion removed as wastewater similar to the LSFO process. A variety

of the lime based wet FGD system is known as a hi-calcium wet lime system. This system does not rely on the

magnesium sulfite inherent in the MEL system as a source of alkalinity. A high calcium wet lime system includes

in-situ forced oxidation and operates in a manner similar to an limestone forced oxidation system. The MEL

system is typically considered to be the more attractive wet lime system of the two options and is therefore the

process evaluated in this study.
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1.1 PROCESS CHEMISTRY

1.1.1 LSFO Process

A simplified explanation of the LSFO process (Figure 1) is that the SO2 absorbed in the recirculated slurry reacts

with dissolved limestone (CaCO3) in the slurry to form calcium sulfite hemihydrate (CaSO3 • 1⁄2H2O) according to

the following simplified reaction:

SO2 + CaCO3 + 1⁄2 H2O ⇒→ CaSO3 • 1⁄2H2O + CO2

Carbon dioxide formed from reaction of limestone with SO2 is released into the flue gas. 

Oxidation air is bubbled through the slurry to convert CaSO3•1⁄2H2O to gypsum (CaSO4•2H2O) according to the

following simplified reaction:

CaSO3 • 1⁄2H2O + 1⁄2O2 + 1.5 H2O ⇒→ CaSO4 • 2H2O

Aeration oxidizes all the calcium sulfite to calcium sulfate and forces precipitation to occur on existing gypsum

crystals in the reaction tank. This minimizes the tendency for gypsum to precipitate on surfaces in the absorber,

piping, and nozzles by maintaining the gypsum concentration in the liquid phase of the absorber slurry liquid at

80-90% of the concentration at which scale would form. The LSFO process also improves slurry dewatering

because the gypsum crystal size and shape dewaters more freely than CaSO3 • 1⁄2H2O crystals produced by older

natural oxidation systems. Physical stability of dewatered gypsum is also improved by the more granular particles,

allowing gypsum to be disposed of in a landfill (if no sales opportunities exist) without requiring fly ash or lime

fixatives.

SO2 collection efficiency and energy consumption of an LSFO system can be improved by adding organic

carboxylic acids to the reaction tank. These additives serve as a buffer to stabilize the pH of the slurry as it
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Figure 1: LSFO Process SO2 Absorber

Recycle Tank

Spray

Headers

Mist

Eliminators

Compressed Air

Limestone

Slurry

Gypsum to

Dewatering

Gas

Inlet

Gas

Outlet

to Stack

M
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absorbs the acidic SO2. This improves absorber performance by increasing solubility of SO2 in the slurry which

allows operation at a lower L/G ratio. It also allows operation at a lower slurry pH, which increases the rate of

dissolution of limestone and increases reagent utilization. (Reagent utilization is the percentage of reagent fed to

the FGD system that reacts with SO2. A portion of the reagent does not react, which increases required reagent

feed rate.) The organic acid is not consumed by the absorption process, but replenishment is required to replace

acid lost through gradual oxidative degradation in the absorber and to replace acid that is carried out in moisture

remaining with the gypsum cake and in the wastewater. 

A number of organic acids have been found to be effective. These acids include adipic, glycolic, maleic, acrylic,

and formic acids. Dibasic acid (DBA), a by-product from the manufacture of adipic acid, is usually selected

instead of other acids because of its lower cost (about $1000 per ton). (DBA is a mixture of adipic, glutaric, and

succinic acids.) If wastewater treatment is required, organic acids have a biological oxygen demand, which may

require removal in an additional water treatment plant. 

In general, LSFO systems are designed without depending on the use of organic acids. Provision is sometimes

included in designs to allow future addition of organic acid in case SO2 collection efficiency is inadequate. 

Disadvantages of organic acids include additional operating cost, uncertainty in long-term supply and pricing of

DBA by-products, possible contamination of gypsum by-product, and increased wastewater treatment cost. 

1.1.2 MEL Process

In the MEL process, slaked lime, containing calcium hydroxide [Ca(OH)2] and a portion of magnesium

hydroxide [Mg(OH)2], is used to react with SO2 (Figure 2). As with the LSFO process, the slurry is added to

a recycle tank at the bottom of the absorber under pH control to replenish reagent consumed. Overall, SO2,

which is a strong acid, reacts with and is neutralized by slaked lime. The overall reactions between lime and

SO2 in the absorber are shown in reactions (1) and (2). Calcium hydroxide in the slurry reacts with most of

the SO2 to precipitate calcium sulfite (CaSO3 • 1⁄2 H2O). Magnesium hydroxide reacts with the remainder of

the SO2 to form soluble magnesium salts; magnesium sulfite and magnesium bisulfite [MgSO3, Mg(HSO3)2]. 
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Figure 2: MEL Process SO2 Absorber with Forced Oxidation

Slaked

Lime

Mist

Eliminators

Spray

Headers

Gas

Inlet

Gas

Outlet

to Stack

Compressed Air

Gypsum to

Dewatering

External

Oxidizer

Recycle Tank

Scoop

M
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These soluble magnesium salts greatly enhance SO2 capture and facilitate the design of a system with lower 

power consumption and equipment costs. When magnesium sulfite is present in slurry in contact with flue

gas, it buffers the pH of the slurry as it absorbs acidic SO2. This improves absorber performance by increasing

solubility of SO2 in the slurry, which can allow operation at a lower L/G ratio. It also allows operation at a lower

slurry pH, near 6.0, which improves reagent utilization to near 100%. Magnesium sulfite is not consumed by the

absorption process, but is replenished by addition of fresh lime slurry to the reaction tank. This provides the

MEL FGD process with a greater capacity to absorb SO2 per specific quantity of absorbing liquor than LSFO.

 Since magnesium sulfite performs the same function as described for organic acids in LSFO systems, no organic

acid addition is required to improve SO2 collection efficiency. A more detailed description of MEL process

chemistry is given in Appendix B. 

The magnesium salts also prevent formation of build-ups on surfaces in the absorber. The salts suppress the

solubility of hard-scale-forming substances like calcium sulfate. This eliminates tendency for calcium sulfate to

precipitate on surfaces in the absorber and cause plugging of pipes and nozzles by maintaining the concentration

of calcium sulfate in the absorber slurry liquid at about 10% of the concentration at which scale would form. 

(1) SO2 + Ca(OH)2 → CaSO3 • 1⁄2 H2O + 1⁄2 H2O

(2) 4SO2 + 3Mg(OH)2 → 2MgSO3 + Mg(HSO3)2 +2H2O

A portion of the recirculated slurry containing absorbed SO2 in the form of calcium sulfite and magnesium

salts is collected in a scoop from the recirculating slurry that has just contacted the flue gas (Figure 2). Slurry
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collected in the scoop has lower pH (~5.5), which makes it more suitable for the subsequent oxidation step. 

The slurry is pumped to an external forced oxidation tank where it is contacted with air. Calcium sulfite is

converted to gypsum and magnesium salts are oxidized to MgSO4, according to reactions (3) and (4) & (5),

respectively. The gypsum precipitates in a crystalline form while MgSO4 remains in solution. 

The gypsum in the slurry effluent from the oxidizer undergoes primary and secondary dewatering to produce

gypsum by-product cake in the same way as in the LSFO process. Liquid containing MgSO4 is returned to

the absorber after dewatering the gypsum. 

(3) CaSO3 • 1⁄2H2O + 1⁄2O2 + 3/2 H2O → CaSO4 • 2H2O

(4) MgSO3 + 1⁄2O2 → MgSO4

(5) Mg(HSO3)2 + O2 → MgSO4 + H2SO4

1.2 REAGENTS AND BY-PRODUCTS

1.2.1 LSFO Process

Preparation of limestone slurry reagent involves grinding a high-calcium, crushed (less than 1 inch) limestone to a

fine size [typically 95% smaller than 325 mesh (44 microns)] in a wet ball mill. The fine size provides for more

complete use of the limestone to maximize reagent utilization and to minimize the amount of unreacted limestone

in the gypsum product. Some LSFO processes claim to be able to accommodate a coarse grind due to a lower

operating pH in the absorber. The ground limestone is slurried with water and held in an agitated tank for use. The

slurry reagent is fed to the absorber to replenish limestone consumed in the reaction and the feed rate is typically

controlled based on pH of the circulating slurry with feed-forward inlet SO2 and flue gas flow rate signals.

The by-product gypsum, after dewatering to 90% solids (10% moisture), can be sold or disposed in a landfill,

either alone or blended with fly ash. Gypsum can be used for making wallboard, as an additive for concrete, or as

an agricultural soil conditioner.
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1.2.2 MEL Process

Lime used in the process is composed of 83 to 91 weight percent calcium oxide (CaO) and 2 to 8 weight

percent magnesium oxide (MgO). Total oxides content is equivalent in SO2 neutralizing value to about 94%

CaO. Magnesium-enhanced lime can be produced from limestone with suitable magnesium content or by

blending high-calcium and dolomitic (~40% MgO) limes. Magnesium oxide content of lime is adjusted by

the supplier to match FGD systems’ requirements for SO2 collection efficiency. This includes adjusting MgO

content so the amount added to the FGD system exceeds the equivalent amount of hydrogen chloride

absorbed. (For example, for 3% sulfur and 0.12% chlorine in coal, 1.2% MgO in lime is required to

accommodate hydrogen chloride; for coal with 1.3% sulfur and 0.1% chlorine, 2.3% MgO is required.) 

Magnesium oxide which reacts with hydrogen chloride is not available to react with SO2 to form magnesium

sulfite, which has the beneficial effects on SO2 collection efficiency and scale prevention discussed earlier. 

To prepare it for use, the lime is mixed with water in a slaker where it is hydrated to form calcium hydroxide

[Ca(OH)2] and magnesium hydroxide [Mg(OH)2]. The resulting slurry is added to the absorber recycle tank

to replenish lime consumed in the process; pH is controlled to a set point near 6.0. At this pH nearly 100% of

the lime reacts with SO2. 

As in the LSFO process, the by-product gypsum, after dewatering to 10% moisture, can be sold or disposed of in a

landfill, either alone or blended with fly ash. This gypsum can be used for manufacturing wallboard, as an

additive for concrete, or as an agricultural soil conditioner. FGD wastewater containing magnesium sulfate and

other soluble species that are undesirable in the gypsum product or in the recirculated scrubber slurry are directed

to a wastewater treatment plant similar to that described for the LSFO process. In this study, wastewater

treatment requirements are assumed to be the same for the MEL and the LSFO processes. Additional cost is

assumed for disposing magnesium sulfate processed through the waste water treatment system. Optionally, a

second by-product, magnesium hydroxide, can be produced by treating the FGD wastewater with additional lime

prior to directing it to the wastewater treatment plant. Magnesium hydroxide has a number of commercial uses,

including injection in the flue gas in the furnace, at the economizer outlet or at the ESP inlet for neutralization of

sulfur trioxide (SO3) in flue gases thereby reducing corrosion in the air heater, and lowering the potential for a

visible plume caused by sulfuric acid condensation. 
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1.3 SO2 REMOVAL PERFORMANCE

1.3.1 LSFO Process

LSFO systems have achieved SO2 removal efficiencies as high as 98% in power plants firing a variety of

high- and low-sulfur fuels. Recent contracts for LSFO technology in the U.S. market have included

guarantees of 99%. These projects are scheduled for start up in 2008 and beyond. For the LSFO absorber, the

higher removal efficiencies are typically attained using a combination of higher L/G’s, higher velocity,

gas/liquid contactors, or through the use of additives such as DBA in the recirculating slurry . This study

employs a design that is not dependent on DBA to achieve the 98% removal requirement.

FGD systems installed in Phase 1 of the CAAA Title IV program (1995) were designed for and achieved 95%

efficiency with L/Gs of 90-130 and inlet sulfur dioxide up to 8 lb/MBtu. Demonstrations and testing by the

FGD process suppliers, including Alstom, Hitachi, Babcock Power, and Wheelabrator, have shown that an

L/G of 130 is adequate to achieve 98% efficiency in a typical open-spray tower design for the high-sulfur coal

case (4.72 lb SO2/MBtu inlet) and an L/G of 80 is appropriate for the low sulfur coal case (2.00 lb SO2/MBtu

inlet). This is further verified by recent guarantees offered by FGD vendors for new and retrofit unit

applications on similar sulfur level coals.

This being the case, for the Appalachian high- and low-sulfur coals, we have estimated that L/G ratios of 130 and

80, respectively, will be required to achieve 98% SO2 removal efficiency without the use of organic acid in a

typical open spray tower LSFO absorber design. Another important design parameter is that limestone reagent

utilization of at least 97% will be necessary to achieve the gypsum quality requirements. 

1.3.2 MEL Process

MEL forced oxidation systems have achieved a better level of performance than the LSFO process, with SO2

removal efficiencies between 98% to 99% in power plants also firing a variety of high- and low-sulfur coals. 

Because of the higher SO2 absorption capacity available in the magnesium-enhanced lime system compared with
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the LSFO system, we estimate that L/G ratios of 40 and 30 will be required to achieve 98% SO2 removal

efficiency for the Appalachian high- and low-sulfur coals, respectively.

The difference in the L/G ratio requirements between the LSFO and MEL processes has a major impact both on

capital and O&M cost differences between the two competing technologies.

1.4 COMMERCIAL STATUS

Both the LSFO and the MEL FGD systems are operating successfully at many coal-fired power facilities, ranging

in size from less than 100 MW to 1000 MW and for coal sulfur content well above 5%. Many competing designs

are available in the marketplace, from a number of well-respected and viable suppliers, including, but not limited

to, the following:

• Alstom

• Babcock & Wilcox

• Babcock Power, Inc.

• Chiyoda

• Marsulex

• Mitsubishi

• Hitachi

• Wheelabrator Air Pollution Control (Siemens)

Although these competing designs vary significantly in their use of the typical mass transfer concepts for

achieving high SO2 removal required of modern FGD systems, recent competition in the U.S. market has shown

that the various designs are very similar in life cycle cost (balance of capital and operating costs). Therefore, for

purposes of this study, we have chosen a conventional open spray tower design (currently offered by at least five

suppliers) as the basis for this work.
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2. WET FGD TECHNOLOGY COMPARISON WITH DRY FGD TECHNOLOGY

2.1 PROCESS ADVANTAGES

The LSFO and MEL technologies have the following advantages when compared with other FGD technologies:

1. Well-established FGD technology on a variety of world coals with proven reliability.

2. SO2 removals of 95% are common and removals as high as 99% have been guaranteed in the

last two years.

3. Multiple and commercially viable suppliers offer the technology.

4. Reagents used by the process are plentiful and readily available.

5. Waste gypsum is stable for landfill disposal without blending with fly ash and lime.

6. Can be designed to produce wallboard-grade gypsum as a saleable by-product.

7. The FGD system is not sensitive to boiler operational upsets and typical operating modes,

such as cycling duty.

2.2 PROCESS DISADVANTAGES

The LSFO and MEL technologies can have the following disadvantages when compared with other FGD

technologies:

1. The LSFO process circulates large quantities of slurry with the attendant high pumping

power consumption.

2. The pressure drop across the absorber increases the induced draft (ID) fan power

consumption.

3. These processes can produce a large volume of gypsum. The salability of this by-product is

dependent on a access to sufficiently sized gypsum market.

4. The high potential for corrosion requires extensive use of costly corrosion-resistant alloys or

nonmetallic liners as materials of construction for the absorber and other system components.
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5. In most applications, a waste water treatment facility will be required to remove heavy metals

from the blowdown stream.
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3. DESIGN BASIS

3.1 SPECIFIC DESIGN CRITERIA

Table 3.1 -1 lists the specific design criteria.

TABLE 3.1-1

SPECIFIC DESIGN CRITERIA

LSFO LSFO MEL MEL

Unit capacity 500 MW 500 MW 500 MW 500 MW

Heat input to boiler, MBtu/hr 5000 5000 5000 5000

Fuel High-sulfur 

Appalachian

Low-sulfur -

Appalachian

High-sulfur -

Appalachian

Low-sulfur -

Appalachian

Fuel analysis, % wt.:

Moisture 3.1 6.0 3.1 6.0

Ash 9.0 9.1 9.0 9.1

Carbon 69.82 72.6 69.82 72.6

Hydrogen 5.0 4.8 5.0 4.8

Nitrogen 1.26 1.4 1.26 1.4

Sulfur 3.0 1.3 3.0 1.3

Oxygen 8.7 4.7 8.7 4.7

Chlorine 0.12 0.1 0.12 0.1

Higher heating value, Btu/lb 12,720 13,100 12,720 13,100

SO2 generation*, lb/MBtu 4.72 2.0 4.72 2.0

Coal feed rate, tons/hr 197 191 197 191

Flue gas flow at FGD inlet**, macfm 1.75 1.70 1.75 1.70

Flue gas temperature at FGD inlet, °F 300 280 300 280

Flue gas flow at FGD outlet, macfm 1.52 1.50 1.52 1.50
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TABLE 3.1-1

SPECIFIC DESIGN CRITERIA

LSFO LSFO MEL MEL

Flue gas temperature at FGD outlet,

°F

130 130 130 130

SO2 reduction efficiency, % 98 98 98 98

SO2 outlet, lb/MBtu 0.094 0.04 0.094 0.04

Mercury concentration in coal, ppmd 0.08 to 0.15 0.08 to 0.15 0.08 to 0.15 0.08 to 0.15

* For material balance purposes, all sulfur in coal is assumed to be converted to SO2 (no SO3) and there is no

sulfur removed in the either the bottom ash or electrostatic precipitator ash.

** The air preheater leakage for the new plant and retrofit plants are assumed to be the same over a long

duration.

Table 3.1-2 summarizes the parameters used for the FGD comparison.

TABLE 3.1-2

PARAMETERS USED FOR FGD COMPARISON

LSFO LSFO MEL MEL

Unit capacity 500 MW 500 MW 500 MW 500 MW

Heat input to boiler, MBtu/hr 5000 5000 5000 5000

Fuel High-sulfur -

Appalachian

Low-sulfur -

Appalachian

High-sulfur -

Appalachian

Low-sulfur -

Appalachian

SO2 removal, % 98% 98% 98% 98%

SO2 emission, lb/MBtu 0.094 0.04 0.094 0.04

By product Gypsum Gypsum Gypsum Gypsum

Power consumption*, % of MWgross 2.10 1.4 1.3 1.0

Reagent Limestone Limestone MEL MEL

Reagent cost, $/ton 15 15 65 65

Reagent purity, % 95 95 94 94

L/G ratio, gpm per 1000 acfm at

absorber outlet

130 80 40 30

Reagent stoichiometry**, moles of

reagent/mole of sulfur removal

1.03 1.03 1.02 1.02
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TABLE 3.1-2

PARAMETERS USED FOR FGD COMPARISON

LSFO LSFO MEL MEL

SO2 oxidation stoichiometry (O/SO2

removed)

3.0 3.0 3.0 3.0

Load factor, % 80 80 80 80

FGD system life, years:

New unit application 30 30 30 30

Retrofit application 20 20 20 20

Capital cost leveling factor, %/year:

New unit application 14.8 14.8 14.8 14.8

Retrofit application 15.7 15.7 15.7 15.7

Discount rate, % 9.0 9.0 9.0 9.0

Inflation rate, % 3.0 3.0 3.0 3.0

Operating cost levelization factor:

New unit application 1.365 1.365 1.365 1.365

Retrofit application 1.274 1.274 1.274 1.274

*Auxiliary power, including estimated calculations, is summarized in Appendix A.

** If the LSFO process is to produce a saleable grade gypsum, a reagent stoichiometry (moles reagent per

mole SO2 absorbed, equal to inverse of reagent utilization) of 1.03 or lower is required, along with greater

than 95% limestone purity. The capital cost takes into account the size of the reaction tank to achieve higher

reagent utilization and crystal growth for the LSFO process.

3.2 SYSTEM COMPONENTS

The design criteria and features selected both for the LSFO and the MEL FGD systems are considered typical for

today’s standard practices. The following descriptions present important features of the components and

assumptions used in this study:
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3.2.1 Reagent Handling and Preparation Systems

3.2.1.1 LSFO Process

Limestone is received by truck and stored in a 30-day capacity bulk storage limestone pile. The reclaim system

includes a yard hopper with a vibrating feeder and conveyor system to transfer limestone to a 24-hour capacity day

bin. The limestone day bin and a gravimetric feeder supply the limestone to a 2 x 100% capacity horizontal ball

mill system. This will allow 100% availability for the limestone grinding system. The ball mill grinds the

limestone to 90% less than 325 mesh and uses a wet recycle classification loop to ensure proper size distribution in

the product. Two 100% capacity classification pumps are used to provide high reliability of the classification

system. The process makeup water or the recycle water is added to the ball mills to produce 70% solid slurry. 

The slurry is diluted to 30% solids in the classification process and is stored in a 16-hour limestone slurry tank

prior to being fed as reagent makeup into an absorber.

3.2.1.2 MEL Process

Lime is received by truck and stored in a 14-day capacity bulk storage lime silo. The lime is pneumatically

conveyed to a 24-hour capacity day bin. The lime day bin and a gravimetric feeder supply the lime to a 2 x 100%

capacity horizontal ball mill slaking system. This will allow 100% availability for the slaking system. A modern,

conventional horizontal ball mill slaker is used. Two 100% slurry transfer pumps are used to provide high

reliability to transfer the slurry to the 16-hour slurry tank. Process makeup water is added to the ball mill slaker to

produce 20% solids product slurry. The slurry is fed to the absorber by a dedicated reagent feed pump (100%

spare capacity provided).

A 100-gallon capacity sulfuric acid tank with sulfuric acid feed pumps is also provided.

3.2.2 SO2 Removal System

3.2.2.1 LSFO Process

One 100% capacity absorber is provided to achieve 98% SO2 removal efficiency. The absorber is a vertical openspray tower, with countercurrent contact between the flue gas and scrubber slurry. To achieve 98% SO2 reduction

with high-sulfur coal, 4 operating pumps will be provided to achieve a 130 L/G ratio. Three operating pumps will
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be provided for low-sulfur coal to achieve an 80 L/G ratio. Each individual pump will be dedicated to a spray

level in the absorber. One spare spray level is provided along with a dedicated pump with each design. The

entrained slurry droplets are removed from the flue gas through a chevron-type mist eliminator. The scrubber

slurry drains into a reaction tank located in the bottom of the absorber. This tank is agitated and is sized for a

minimum of 15 hours of solids residence time. Compressed air is injected below the reaction tank agitator through

a sparger network. To achieve greater than 99% oxidation of CaSO3 to CaSO4, a 3:1 stoichiometric ratio of

oxygen to absorbed SO2 (moles O/moles SO2) is used to size the oxidation air compressors. The bleed slurry from

the absorber is pumped to the dewatering area. Makeup limestone slurry is added to the reaction tank. The

material of construction for the absorber is UNS S32205 steel. The chloride levels will be maintained at less than

15,000 ppm.

The majority of water loss is through evaporation, which is assumed to be the same for both processes (LSFO

and MEL). Since it is also assumed that both processes remove the same amount of sulfur from the gas,

converting it into a saleable grade gypsum, both processes will have the same water losses in the gypsum byproduct. The wastewater blowdown is based on the chloride level, which, again, will be the same for both

processes. Therefore, the water losses and consumption will be the same for both wet processes.

3.2.2.2 MEL Process

One 100% capacity absorber is provided to achieve 98% SO2 removal efficiency. The absorber is a vertical openspray tower, with countercurrent contact between the flue gas and scrubber slurry. To achieve 98% SO2 reduction

with high-sulfur coal, 2 operating pumps will be provided to achieve a 40 L/G ratio. Two operating pumps will

also be provided for low-sulfur coal to achieve a 30 L/G ratio. Each individual pump will serve a dedicated spray

level in the absorber. One spare spray level is provided, along with a dedicated pump, for both L/G designs. The

absorber is constructed of UNS S32205 steel or equivalent. The chloride levels will be maintained at less than

15,000 ppm. The entrained slurry droplets are removed from the flue gas through a chevron mist eliminator. 

Makeup lime slurry is added to the recycle tank. A trough (scoop) system is employed to direct a small portion of

the recirculated slurry to an oxidation tank for the conversion of CaSO3 to CaSO4. The remaining slurry returns to

the recycle tank.
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 The oxidation tank is sized for 4 hours of solids residence time. To achieve greater than 99% oxidation of CaSO3

to CaSO4, a 3:1 stoichiometric ratio of oxygen to removed SO2 is used in the oxidation compressor design. 

Compressed air is injected below the oxidation tank agitator through a sparger network. The bleed slurry from the

oxidation tank is pumped to the dewatering area, where a conventional gypsum dewatering system, similar to the

LSFO process, is employed. 

As an alternative to the conventional dewatering system, S&L contacted EIMCO to discuss the application of the

OxallizersTM process to the hypothetical applications of the MEL process in this study. On a total cost of electricity

impact basis, including both capital and operating costs, the OxallizersTM process is approximately 10% more

costly than the conventional process considering the assumptions used in this analysis.

Therefore, the capital costs presented in this report are based on the conventional sparged oxidation system rather

than the OxallizerTM.

3.2.3 Flue Gas System and Stack

3.2.3.1 LSFO Process

The gas path of the system will start at the discharge of the existing ID fans, through the new booster ID fans and

absorber, and discharge into a new FRP lined chimney. The draft requirement of the LSFO FGD system will be

accommodated by the booster ID fans. The booster ID fan is sized to provide an additional 10.4" H2O (9.4"

operating) and 9.2” H2O (8.2" operating) pressure drop through the absorber for high- and low-sulfur coal

applications, respectively. The inlet ductwork to the absorber system, including the booster ID fan, will be

fabricated with carbon steel. The outlet ductwork from the absorber to the stack breeching will be constructed of a

carbon steel substrate clad with a 1/16-inch “layer” of C-276 alloy for corrosion protection.

3.2.3.2 MEL Process

The flue gas system for the MEL process will be identical to the LSFO process, with the exception that the draft

requirements will be reduced to 8.1” H2O (7.1” operating) and 7.8” H2O (6.8” operating) for the high- and lowsulfur cases, respectively.
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3.2.4 Gypsum Dewatering and Handling System

The gypsum dewatering system will be similar for both the LSFO and MEL processes.

The bleed stream from the reaction tanks for the LSFO process and from the oxidation tank for the MEL process is

pumped to a set of hydroclones for primary dewatering. The hydroclone overflow is recycled back to the process.

 The hydroclone underflow, containing 50% solids, is fed to a belt filter for secondary dewatering. Two 100%

horizontal belt filters are designed to achieve 90% solids in the gypsum by-product. The gypsum cake is washed

with fresh water on the belt filters to remove any residual chlorides to a level of less than 100 ppm of chloride in

the dry solids. The filtrate from the horizontal vacuum belt filter is returned to the process. A small blowdown

from the process filtrate is required to remove chloride from the process. For purposes of this study, we have

assumed that this chloride purge will have to be treated in a wastewater treatment facility to remove heavy metals

before discharge.

The dewatered gypsum is discharged from the belt filters to belt conveyors that transfer the gypsum to a covered

storage area. The covered storage area is sized for 7 days. The gypsum is assumed to be trucked to an offsite user.

3.2.5 General Plant Support Systems

The general support equipment includes modifications and/or additions to the plant seal water system, instrument

air system, makeup water system, distributed control system (DCS), and electrical auxiliary power system and

accommodations in the BTG control room. Additionally, a typical FGD facility will require plant site

accommodations for roads, storm sewers, sanitary systems, and fire protection systems. These accommodations

are included in the design and cost estimate.
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Table 3.2.5–1 lists the equipment used in each subsystem.

TABLE 3.2.5 -1

EQUIPMENT USED IN EACH SUBSYSTEM

Reagent Handling and Preparation

LSFO Process:

Truck unloading system

Limestone storage pile (30 days storage)

Limestone reclaim system

Limestone day bins(24 hours storage)

Limestone crushers

Limestone ball mills (2 x 100% capacity) system

Slurry storage tank (16 hours storage)

MEL Process:

Truck unloading system

Lime bulk storage steel silo (14 days storage)

Lime live reclaim system

Lime day bins (24 hours storage)

Lime ball mill slakers (2 x 100%% capacity)

Slurry storage tank (16 hours storage)

SO2 Removal System

LSFO Process:

1 x 100% module with reaction tank and spargers

5 recirculation pumps (4 operating, 1 spare)

2 x 100% air compressors

MEL Process:

1 x 100% module with reaction tank and slurry trough

3 recirculation pumps (2 operating, 1 spare)

Oxidation tank with spargers

2 x 100% air compressors
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By-Product Treatment and Handling System

Hydroclones (20% spare hydroclones)

Hydroclone underflow surge tank

Belt filters (2 x 100%)

Conveyor to storage shed

Gypsum storage shed

Recycle process water tank

Flue Gas System

Booster ID fans (2 x 50%)

Absorber inlet ductwork/dampers

Absorber outlet ductwork

FRP lined stack for retrofit applications

Bypass duct to existing stack with damper

General Plant Support Systems

Makeup water tank, pumps, piping

Instrumentation/plant air compressors (2 x 50%)

Auxiliary power transformers/switchgear/motor controls, etc.

Electrical conduit, duct banks, wiring, cables, etc.

Control system (DCS)

Site accommodations:

Roads

Storm sewers

Sanitary systems

Fire protection
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4. TECHNOLOGY APPLICATION CONSTRAINTS

Summarized below are the typical application constraints for conventional wet FGD technology.

4.1 UNIT/ABSORBER SIZE

Conventional wet FGD systems are in successful operation at a large number of coal-fired facilities ranging in

size from small, individual applications to large utility applications of more than 1000 MW. Single absorbers

capable of treating the flue gas flow from a 1000 MW bituminous coal-fired plant are commercially

demonstrated and available from industry suppliers. Therefore, the design basis for this study of a single

absorber for a 500 MW facility is well within the current range of commercial applications of FGD

technology.

4.2 COAL SULFUR CONTENT

Wet FGD systems have been successfully applied to coals of all sulfur levels. Most of these systems are

installed on medium- to high-sulfur fuels (>2%); however, a number of systems are also installed on lowsulfur fuels (<1%) in the western areas of the United States.

4.3 PERFORMANCE EXPECTATIONS

SO2 removal guarantees of up to 99% are available from the system suppliers and have been demonstrated in

commercial applications.

4.4 REAGENT UTILIZATION

Reagent utilization is the percentage of reagent fed to the FGD system that reacts with SO2. Reagent utilization in

a wet FGD system is typically very close to the required theoretical level when considered on a removed SO2

basis. Limestone reagent utilization in LSFO systems is typically greater than 97%. This high level of

reagent utilization is also required to maintain the quality of gypsum if it is to be used for wallboard
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manufacturing. Lime reagent utilization is near 100% due to the high reactivity of lime and small particle size

of slaked lime (3 microns average). 

4.5 BY-PRODUCT QUALITY

The by-product from either the wet LSFO process or the MEL process can be fashioned for use in the cement

industry, for wallboard manufacturing, or for agricultural use. The by-product quality requirements are much

more stringent for wallboard manufacturing compared to the other uses and will typically require the use of a

gypsum washing system and subsequent water treatment system to handle soluble impurities not acceptable to

wallboard manufacturers.

4.6 ENERGY CONSUMPTION

The process energy consumption results primarily from the booster ID fan power required to overcome the

draft loss across the absorber and the power requirement for recirculation pumps. The MEL process has an

advantage over the LSFO process because the power requirement for slurry recirculation is 74% lower (highsulfur case) than for the LSFO process and the power requirement for the booster ID fan is 27% lower (highsulfur case) than for the LSFO process. Overall, the MEL process will require approximately 42.8% less

auxiliary (% of MWgross) power for high-sulfur coal and approximately 36.9% less auxiliary power auxiliary

(% of MWgross) for low-sulfur coal compared to the LSFO process. 

4.7 RETROFIT VERSUS NEW UNITS

In new unit and retrofit applications, wet FGD technology is typically installed between the electrostatic

precipitator/baghouse outlet and the stack. 

Most of the retrofit units will have booster fans to overcome the pressure drop across the FGD absorber,

which will be located after the existing ID fans. However, new units are expected to be installed with ID fans

large enough to overcome the pressure drop across the FGD absorber. This feature will typically result in a

lower capital cost for the draft system on a new unit application compared with a retrofit application.
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Most retrofit units will not be able to use existing stacks, as these stacks are typically designed for hot flue gas

at approximately 100 ft/sec exit velocity. To accommodate the saturated flue gas from a wet FGD system,

wet stacks are typically designed for a highly corrosion-resistant material with a gas velocity of between 55 to

70 ft/sec. The lower gas velocity is required to prevent condensed moisture from being carried out the top of

the stack. Most of the retrofit units will, therefore, require a new wet stack. 

In some cases, retrofit of the FGD equipment will be hampered by the existing site constraints which may

result in unique and not well optimized equipment arrangements that would not be the case for new units.

Other than these three key areas, an FGD system itself will be very similar for either the retrofit or the new fit

application.

4.8 HG REMOVAL

The 1990 Clean Air Act Amendments required the EPA to study the public health risks, if any, associated

with hazardous air pollutant emissions, including mercury, from fossil fuel-fired power plants. The Utility

Study was completed in the late 1990’s and identified uncertainties with respect to mercury emissions from

coal-fired power plants. In order to address uncertainties identified in the mercury reports, EPA issued an

information collection request (ICR) requiring all coal-fired power plants to collect and analyze coal samples

at their facilities. The ICR also required approximately eighty power plants representing various combinations

of fuel types and control technologies to perform mercury speciation tests on stack emissions. This lack of

data regarding mercury emissions was a great obstacle to formulating approaches to control mercury

emissions from power plants. The ICR was implemented, and a data base of mercury emissions was

developed. The existence of the ICR data initiated R/D efforts and eventually revealed some cost effective

approaches for reducing mercury emissions. 

The ICR data showed that mercury is present in power plant emissions as both solid and vaporous forms. The

solid form is a particulate of a size and character that is collected effectively in the existing particulate control

equipment (fabric filters and precipitators) typically installed on power plants. Hence, some of the mercury in

the coal is already being collected; however, the ICR data also revealed that a majority of the mercury is not

in the solid particulate form; rather the mercury is in a vaporous form. 
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The ICR data indicated that the vaporous form of mercury predominates in power plant exhaust gasses. The

vaporous form was categorized as two species, elemental and ionic mercury. Elemental mercury is a

molecule that is unreacted, neutral valence, and very stable in the flue gas. The data showed that elemental

mercury, Hg0, is only minimally collected in the current power plant emission control equipment. The

second category, or species, of mercury is ionic mercury, Hg+2. Ionic mercury is a compound of mercury

and other molecules and has a positive valence. The data showed that ionic mercury is collected in some of

the existing power plants’ pollution control equipment. 

The ICR data did not offer a clear correlation between mercury collection and the existing control devices,

until researchers identified that power plants firing eastern bituminous coals had greater removal efficiencies

and also tended to have greater concentrations of chlorides in their coal analysis. The discovery of this

correlation has helped define the problem of controlling power plant mercury emissions. The correlation

revealed that chlorides in the coal were associated with higher ionic mercury concentrations. Ionic mercury

can be removed with existing pollution control equipment. Coals with higher coal chloride levels exhibited

significantly higher mercury removal efficiencies at existing power plants. Further study of ionic mercury

compounds supported the correlation, because many of these ionic species were chloride and mercury

compounds. The following table illustrates the general tendency found in the ICR data for the speciation of

mercury in Eastern Bituminous, PRB Sub-Bituminous, and Lignite coals. 

Mercury Speciation in Flue Gas for Various Coal Ranks

Coal Type Median Hg

in Coal

Flue Gas

Particulate

Bound Hg

Flue Gas

Ionic Hg

Flue Gas

Elemental

Hg

Chloride

Content in Coal

Eastern

Bituminous

7 lb/TBtu 44.49% 30.25% 25.26% 500-1000 ppm

PRB SubBituminous5 lb/TBtu 17.56% 17.44% 65.01% 25-200 ppm

Lignite 8 lb/TBtu 17.96% 18.88% 63.16% 25-300 ppm
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This correlation explains many of the differing removal efficiencies exhibited at power plants. Plants that fire

Powder River Basin (PRB) coals have relatively low amounts of chloride in the coal, have greater

concentrations of elemental mercury, and they exhibit low mercury removal efficiencies. Plants firing

bituminous coals have higher chloride concentrations in the coal, lower concentrations of elemental mercury,

and have greater mercury removal efficiencies. 

In support of this correlation, further evaluation of the ICR data identified that baghouses remove

considerably more mercury than precipitators, but that the removal depends on the coal type. The removal for

baghouses collecting PRB fly ash ranged from 53% to 88%; and 67% to 93% when bituminous coals were

fired. It is presumed that baghouses performed better when bituminous coals were fired because they had

more ionic mercury in the flue gas, and the baghouses were collecting more of this ionic mercury. Elemental

mercury is also being collected in the baghouse but not nearly as efficiently as is the ionic. These results were

without any sorbent injection. 

Also supporting the correlation, the ICR data indicated wet FGD systems can remove approximately 80% of

the inlet mercury from bituminous fired boilers, but only 35% removal on PRB fired boiler. Again, it appears

the ionic mercury is collected effectively in the wet FGD system, but the elemental is not collected at all. The

PRB has mostly elemental mercury and showed considerable less mercury capture in the wet FGD system. 

A second correlation that resulted from evaluation of the ICR data showed greater removal of ionic mercury

in the emission reduction equipment that operated at air heater gas outlet temperatures. This was evident when

evaluators identified that ionic mercury was not readily collected in hot side precipitators. While the

temperature of the flue gas affects mercury collection, this correlation is not nearly as important as the

correlation with chlorides. This result has focused investigators efforts at collection at air heater outlet

temperatures. 

While these two correlations resulted from analysis of the ICR data, the data still reveals inconsistencies and

much variation in the removal data. The ICR data dramatically improved the understanding of mercury

removal, but there is still much investigation needed to fully understand mercury speciation and mercury

capture and how coal chemistry is related to each issue.
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The availability of the ICR data resulted in an infusion of grant moneys to evaluate the control of both

elemental and ionic mercury emissions from power plants. In the last six years, R/D has identified many

emission reduction alternatives and characterized their applicability to coal-fired power plants; however, few

of the technologies have yet received long term demonstration in a power plant situation. Thus, there is much

still remaining with regard to knowing how technologies will perform in a specific application and the

reliability of the systems.

In conclusion, the preponderance of the data on Mercury speciation in the combustion process and the

subsequent collection in post combustion emission control technologies indicate that there will be no

significant difference between the LSFO and the MEL FGD processes.
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5. COSTS ANALYSES

5.1 CAPITAL COSTS

Estimated capital costs for the high- and low-sulfur coal applications for both the LSFO and the MEL wet

FGD systems were developed for new and retrofit applications. These estimates encompass all of the

required equipment, materials, and construction associated with the new and retrofit installations of these

technologies. 

The costs were developed primarily from Sargent & Lundy’s cost database, which is continuously updated

from ongoing work in the area of FGD.

The capital cost estimates provided herein represent the “total plant cost,” and include the following:

• Equipment and material (FGD system and balance-of-plant)

• Direct field labor

• Indirect field costs and engineering

• Contingency

• Owner's cost

• Allowance for funds used during construction (AFUDC)

• Initial inventory and spare parts

• Startup and commissioning

The capital cost estimates provided do not encompass sales tax or property tax. License fees and royalties are

not expected for these FGD technologies. Owner’s cost includes owner’s project management, engineering,

accounting, billing, etc., as well as any “corporate” charges that may apply to capital projects. Capital costs

are in 2006 dollars and therefore no escalation is included.

Salient features of each capital cost estimate are described in Section 3 of this report.
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Based on our experience and the level of effort applied to this study, we estimate that the relative accuracy of

the estimate prepared is ±20%.

Additionally, an underlying assumption is that the contracting arrangement for the implementation of an FGD

project is based on multiple lump sum specialty work packages. If the FGD project is to be executed on an

engineer, procure, construct (EPC) or turnkey basis, a separate risk allocation should be added to the estimate

of 10% to 20% (1.1 or 1.2 multiplier) to cover the EPC contractor’s costs, contingency, subcontract mark-ups,

and project risk profile.

Exhibit 5–1 and Exhibit 5–2 present the capital costs for new units and retrofit units, respectively.

5.2 OPERATIONS AND MAINTENANCE COSTS

Exhibit 5–3 through Exhibit 5–6 present the estimated O&M costs associated with the wet and dry FGD

systems. These include both fixed and variable operating costs, defined as follows:

5.2.1 Fixed O&M Costs

The fixed O&M costs consist of O&M labor, maintenance material, and administrative labor. 

For purposes of this study, the installation of the FGD system has been anticipated to add an additional 8

operators to the current pool of operating labor for new units and 12 operators for the retrofit application. For

retrofit applications, we have assumed that the plant layout and integration between the existing power

generation block and the FGD system is not fully optimized, which would require more operating labor than

would a new unit. Four extra personnel are added for a retrofit application to cover the impact of a non-laborefficient plant layout and control room. 

Maintenance material and labor costs shown herein have been estimated based on technology operating

experience in the United States and Europe. The maintenance cost includes periodic replacement of

equipment, the required maintenance material, and the labor required to perform the maintenance.
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5.2.2 Variable O&M Costs

Variable O&M costs developed for each FGD system include the cost of all consumables, including reagent

(limestone or lime), water, and power requirements. In each case, we have assumed that gypsum will be

“sold” for $0/ton.

Auxiliary power costs reflect the power requirements associated with the operation of the fans (incremental

ID fan power for new units and booster ID fan power for retrofit units) as well as the power consumption for

the FGD system, reagent handling, reagent preparation, gypsum dewatering, gypsum handling, and

miscellaneous ball mills/slakers, recirculation slurry pumps, compressor for oxidation air, gypsum dewatering

system, and electrical loads and control users typically needed for FGD operations.

Considering both the regulated and unregulated market and new and retrofit applications, Sargent & Lundy

believes that $45/MWh is representative of the current power market. Auxiliary power cost for retrofit

applications tend toward the lower end of the range; whereas, for new applications, these costs tend toward

the higher end of the range.

Exhibit 5–3 and Exhibit 5–4 represent the fixed and variable O&M costs for new and retrofit applications of

the LSFO FGD technology, respectively. Exhibit 5-5 and Exhibit 5-6 represent the fixed and variable O&M

costs for new and retrofit applications of the MEL FGD technology, respectively. Included with each of these

Exhibits are the assumptions used for all the components in the fixed and variable O&M costs.

5.3 LEVELIZED COSTS

Levelized costs, also referred to as “life cycle costs,” combine the one time capital costs and the yearly O&M

cost over the life of the plant. The levelized fixed charge rate (impact due to capital cost) was calculated

based on an assumption that a plant owner is a regulated utility. The levelized fixed charge rate includes

depreciation of the property, return on capital (50% debt and 50% equity), income tax, property tax, and

insurance. Based on a 9.0% discount rate and 30-year or 20-year life expectancy for new or retrofit facilities,

respectively, the levelized fixed charge rates are 14.8% (30-year life) and 15.7% (20-year life). The levelized
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cost analysis was performed based on current dollars, as most regulated utilities base their analysis on current

dollars.

The levelized O&M cost factor takes into account the discount rate, escalation rate, and annuity rate. The

levelized O&M cost factors are 1.365 for a 30-year analysis and 1.274 for a 20-year analysis.

Exhibits 5-7 through 5-10 summarize the levelized costs for the four cases of this study; LSFO on new and

retrofit units and MEL on new and retrofit units, respectively. A summary of the levelized costs in ¢/kW-hr

for each of the study cases for both the high- and low-sulfur coals is as follows:

Levelized Costs (¢/kW-hr)

New Units

High Sulfur Coal Low Sulfur Coal

LSFO 0.80 0.66

MEL 0.82 0.64

Retrofit Units

High Sulfur Coal Low Sulfur Coal

LSFO 0.99 0.85

MEL 1.00 0.82
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APPENDIX A

AUXILIARY POWER SUMMARY

One of the major components of auxiliary power is additional power required by ID fans to overcome system

pressure drop caused by flue gas flow through the absorber. The system pressure drop was estimated as

follows in Table A-1:

TABLE A-1

ESTIMATED FGD SYSTEM PRESSURE DROP

LSFO HighSulfurLSFO LowSulfurMEL HighSulfurMEL LowSulfur

Ductwork to absorber

(“w.c.)*

11 1 1

Absorber inlet expansion

(“w.c.)

0.5 0.5 0.5 0.5

Absorber spray (“w.c.) 3.64 2.24 1.12 0.84

L/G 130 80 40 30

“w.c. loss per 25 L/G

(assumed)

0.7 0.7 0.7 0.7

Mist eliminator (“w.c.) 0.8 0.8 0.8 0.8

Absorber outlet contraction

(“w.c.)

0.5 0.5 0.5 0.5

Ductwork to stack (“w.c.) 0.2 0.2 0.2 0.2

Additional stack pressure due

to wet stack (“w.c.)

22 2 2

Margin (“w.c.) 1 1 1 1

Total 9.64 8.24 7.12 6.84

*”w.c.=pressure loss in

inches of water column

The power requirements with flue gas handling, based on the above pressure drop, and the power requirement

for other FGD subsystems were calculated as follows in Table A-2:
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TABLE A-2

POWER REQUIREMENTS WITH FLUE GAS HANDLING AND OTHER FGD SUBSYSTEMS

LSFO HighSulfurLSFO LowSulfurMEL HighSulfurMEL LowSulfur

Reagent Preparation (kW) 720 305 450 190

SO2 Absorption (kW) 4,463 2,692 1,161 859

No. operating spray

levels

43 2 2

#4 (highest) header

elevation (ft)

70.0 -- -- --

#3 header elevation

(ft)

65.0 65.0 -- --

#2 header elevation

(ft)

60.0 60.0 50.0 50.0

#1 (lowest) header

elevation (ft)

55.0 55.0 45.0 45.0

Average header

elevation (ft)

62.5 60.0 47.5 47.5

Nozzle pressure loss

(psi)

10 10 10 10

Specific gravity of

pumped slurry

1.09 1.09 1.16 1.16

GPM slurry 197,600 120,615 60,800 45,000

Pump efficiency (%) 85.00 85.00 85.00 85.00

Motor efficiency (%) 95 95 95 95

Oxidation Air Compression

(kW)

1,350 572 1,350 572

Flue Gas Handling (kW) 2,592 2,216 1,915 1,839

By-product Handling (kW) 420 178 420 178

General (kW) 60 60 60 60

Total (kW) 9,440 5,865 5,355 3,699

Parasitic load (% of

generation)

1.89 1.17 1.07 0.74
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Margin (%) 0.2 0.2 0.2 0.2

Total (%) 2.09 1.37 1.27 0.94

Parasitic load used in

study(%)

2.1 1.40 1.30 1.00

The booster fan and slurry pump efficiencies used in the study are 85% and the associated motor drives are

95%, respectively. The LSFO slurry pump curves at 50,000 gpm show a somewhat higher pump efficiency of

87-89%, whereas the MEL pump curves at 30,000 gpm show a lower efficiency of 81-83% in the area of

required developed head. It is assumed, for the purposes of this study, that the same efficiency pumps will be

available for both services.
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APPENDIX B

DETAILED DESCRIPTION OF MEL PROCESS CHEMISTRY

Steps in Absorption of SO2 in MEL Process

In the MEL process, slaked lime, containing calcium hydroxide [Ca(OH)2] and a portion of magnesium

hydroxide [Mg(OH)2], is used to react with SO2. Slaked lime is added to the recycle tank of an SO2 absorber

to replenish reagent consumed, as shown in Figure 1. 

Figure 1: Reactions in MEL SO2 Absorber

SO + H O HSO + H

HSO + H + MgSO Mg(HSO )

22 3

3 3 32

- +

- +

Mg(HSO ) + Ca(OH) 

CaSO 1⁄2 H O (solid) + 

MgSO (aqueous)+ 1-1/2 H O

2 Mg(HSO ) + Mg(OH) 

2 MgSO + + 2 H O

32 2

3 2

3 2

32 2

3 2 Mg(HSO )3 2 

102 SO + 94 Ca(OH) + 6 Mg(OH)

94 CaSO 1⁄2 H O (solid) + 

4 MgSO (aq.) + 2 

2 2 2 

3 2

3

 51 H O

Mg(HSO ) (aq.)

2

3 2 

STEP 1: ABSORPTION

STEP 2: PRECIPITATION

AND REGENERATION:

OVERALL:

SO2

SO ABSORBED

BY MgSO

2

3

Mg(HSO )3 2

MgSO in solution 3

Flue gas

with SO2

Cleaned

Gas

To Forced

Oxidation

& Gypsum

Production

Slaked Lime

94 moles

Ca(OH)

6 moles

Mg(OH)

2

2

CaSO 1⁄2 H O 3 2 M

The overall reaction in the absorber of lime with SO2, for a magnesium-enhanced lime with 6 moles of

magnesium hydroxide and 94 moles of calcium hydroxide, is shown in reaction (1):
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(1) 102 SO2 + 94 Ca(OH)2 + 6 Mg(OH)2 → 94 CaSO3 • 1⁄2 H2O

(solid) + 4 MgSO3 (aqueous) + 2 Mg(HSO3)2 (aqueous)

+ 51 H2O

Although reaction (1) shows the overall reaction, SO2 does not react directly with calcium or magnesium

hydroxide. SO2 capture with lime occurs in two steps. In the first step, SO2 is absorbed into the liquid phase

of slurry recirculated from the recycle tank through banks of spray nozzles. In the second step, SO2 absorbed

into the slurry and in the form of an intermediate reaction product falls into the recycle tank where it reacts

with slaked lime. 

In the first step, SO2 is absorbed by recirculated slurry whose liquid phase contains magnesium sulfite

(MgSO3). Magnesium sulfite is alkaline with respect to SO2, which is highly acidic. When SO2 is absorbed

into the liquid, it combines with water to form sulfurous acid, which dissociates into hydrogen and hydrogen

sulfite ions (HSO3

-

), shown in reaction (2). MgSO3 reacts with the hydrogen ions, shown in reaction (3),

which buffers pH so that slurry pH falls less than 0.5 pH units from the recycle tank set point of 6. This

buffering prevents accumulation of SO2 in the liquid by keeping reaction (1) driven far to the right. By

preventing accumulation of SO2 in the liquid, maximum driving force for mass transfer of SO2 from flue gas

to the liquid is maintained. 

(2) SO2 + H2O → H+

 + HSO3

-

(3) H+

 + HSO3

-

 + MgSO3 → Mg(HSO3)2

In the second step, slurry containing captured SO2 in the form of Mg(HSO3)2 falls into the recycle tank. Its

lower pH depresses slightly the pH of the contents of the recycle tank, and lime slurry containing both

calcium and magnesium hydroxides is added to restore pH to the set point. Calcium hydroxide reacts with

most of the Mg(HSO3)2 to precipitate calcium sulfite hemihydrate (CaSO3 • 1⁄2H2O) and regenerate MgSO3,

according to reaction (4). Magnesium hydroxide in the slaked lime reacts with a smaller portion of
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Mg(HSO3)2 to form additional MgSO3 as shown in reaction (5). A portion of captured SO2 remains in the

form of Mg(HSO3)2. 

(4) Mg(HSO3)2 + Ca(OH)2 → CaSO3 • 1⁄2H2O (solid)

+ MgSO3 (aqueous) + 11⁄2H2O

(5) 2Mg(HSO3)2 + Mg(OH)2 → 2MgSO3 (aqueous)

+ Mg(HSO3)2 (aqueous) + 2H2O

A portion of the MgSO3 formed in reactions (4) and (5) is converted to MgSO4 by oxygen present in the flue

gas. MgSO4 does not buffer pH in a range effective for SO2 removal. 

However, a mechanism called co-precipitation allows a portion of this MgSO4 to be converted back to

MgSO3. When CaSO3 • 1⁄2H2O precipitates, it tends to include sulfate (SO4

=

) ions in its crystal lattice in place

of some sulfite (SO3

=

) ions. (This co-precipitation occurs even when the scrubber liquor is unsaturated with

gypsum). As a result, magnesium ions which were associated with the SO4

=

 in MgSO4 become available to

pair with SO3

=

 and regenerate additional MgSO3. Rewriting equation (4) to include this effect, and assuming

that the mole fraction of CaSO4 in the calcium sulfite crystal lattice is 0.2, shows how this additional MgSO3

is formed:

(6) Ca(OH)2 + Mg(HSO3)2 + 0.2 MgSO4 → Ca(SO3)0.8 (SO4)0.2 • 1⁄2H2O (co-precipitate) +

1.2 MgSO3 + H2O

As a result of this effect, a substantial amount of MgSO3 is produced in addition to that produced by addition

of Mg(OH)2 in slaked lime. 

Forced Oxidation and Production of Gypsum

A portion of the recirculated absorber slurry that has just contacted the flue gas is collected in a scoop, located

above the level of slurry in the recycle tank (Figure 1). Slurry is withdrawn from this location because it has a

lower pH which is more suitable for the subsequent oxidation step. Slurry withdrawn from the absorber via
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the scoop contains calcium sulfite solid and magnesium sulfite and magnesium bisulfite in the liquid; the

combined molar rate of withdrawal is equal to the molar absorption rate of SO2. 

The slurry is pumped to an external forced oxidation tank where it is contacted with compressed air. 

CaSO3•1⁄2H2O is converted to gypsum (CaSO4•2H2O), and MgSO3 is oxidized to MgSO4, according to

reactions (7) and (8), respectively. The gypsum precipitates while MgSO4 remains in solution. Liquid

containing MgSO4 is returned to the absorber after dewatering the gypsum. A portion of MgSO4 returned to

the absorber replenishes MgSO3 as described for reaction (6). 

Mg(HSO3)2 in the liquor is oxidized to form MgSO4 and sulfuric acid (H2SO4), shown in reaction (9), which

reduces pH to below 5.5 necessary to speed oxidation of CaSO3•1⁄2H2O. This “free” sulfuric acid replaces

purchased sulfuric acid that would otherwise be required to reduce pH in the external oxidizer. Sulfuric acid

increases dissolution rate of calcium sulfite, which increases overall oxidation rate. It also neutralizes small

amounts of unreacted lime and calcium carbonate carried over from the absorber, shown in reaction (10). As

a result, a negligible amount of unreacted lime remains, which allows overall lime reagent utilization to be

more than 99.9%.

(7) CaSO3 • 1⁄2H2O + 1⁄2O2 + 3/2 H2O → CaSO4 • 2H2O (gypsum)

(8) MgSO3 + 1⁄2 O2 → MgSO4

(9) Mg(HSO3)2 + O2 → MgSO4 + H2SO4 (sulfuric acid)

(10) CaO + CaCO3 +2H2SO4 → 2 CaSO4• 2H2O + CO2
National Lime Association 

Exhibit 5-1: Capital Cost for 500 MW New Unit using Appalachian Coals - Wet FGD 

Limestone Forced Oxidation Magnes

Appalachian High Sulfur Appalachian Low Sulfur Appalachian High Sulfur 

Subsystems Cost, US$ $/kW Cost, US$ $/kW Cost, US$ $/k

 Reagent Feed System 12,825,000 25.7 10,432,800 20.9 8,740,000 17 SO2 Removal System 28,475,000 57.0 25,925,000 51.9 25,627,500 5 Flue Gas System 11,500,000 23.0 11,270,000 22.5 11,250,000 22 By-product Treatment and Handling 9,975,000 20.0 9,025,000 18.1 9,975,000 20 General Support Equipment 5,000,000 10.0 4,500,000 9.0 5,000,000 10 Including Miscellaneous Equipment (Additional  Transformer, Switchgear, wastewater treatment)  TOTAL PROCESS CAPITAL (TPC) 67,775,000 136 61,152,800 122 60,592,500 1 General Facilities (5% of TPC) 3,389,000 6.8 3,058,000 6.1 3,030,000 6 Engineering and Construction Management 6,778,000 13.6 6,115,000 12.2 6,778,000 13 Project Contingency (15%) 11,691,000 23.4 10,549,000 21.1 10,560,000 2 TOTAL PLANT COST (TPLC) 89,633,000 179.3 80,874,800 161.7 80,960,500 16 Allowance for Funds (AFUDC - 7.0% of TPLC) 6,274,000 12.5 5,661,000 11.3 5,667,000 1 Owner's Cost (5% of TPLC) 4,482,000 9.0 4,044,000 8.0 4,048,000 8 TOTAL PLANT INVESTMENT (TPI) 100,389,000 200.8 90,579,800 181.1 90,675,500 18 Inventory Capital (Spare, 1% of TPI) 1,004,000 2.0 906,000 1.8 907,000  Initial Chemicals and Commissioning (2% of TPI) 4.0 1,812,000 2,008,000 3.6 1,814,000 3

 TOTAL CAPITAL REQUIREMENT (TCR) 103,401,000 207 93,297,800 187 93,396,500 1

TOTAL CAPITAL REQUIREMENT (TCR) with BAGHOUSE 128,401,000 257 118,297,800 237 118,396,500 2Notes:1.0 Accuracy of Estimate +/-20%2.0 Labor cost based on normal 5 x 8 shift operation3.0 Project implementation based on multiple lump sum specialty contracting4.0 Costs are based on year 2006 (No escalation)5.0 Baghouse is included in the cost for comparison with dry FGD

New capital/CAPCOST-wet FGD-2006.xls.xls
National Lime Association 

Exhibit 5-2: Capital Cost for 500 MW Retrofit Unit using Appalachian Coals - Wet FGD

Limestone Forced Oxidation Magnesiu

Appalachian High Sulfur Appalachian Low Sulfur Appalachian High Sulfur 

Subsystems Cost, US$ $/kW Cost, US$ $/kW Cost, US$ $/kW 

 Reagent Feed System 13,500,000 27.0 11,340,000 22.7 9,200,000 18.4  SO2 Removal System 33,500,000 67.0 30,500,000 61.0 30,150,000 60.3  Flue Gas System with new stack 23,000,000 46.0 22,540,000 45.1 22,500,000 45.0  By-product Treatment and Handling 10,500,000 21.0 9,500,000 19.0 10,500,000 21.0  General Support Equipment 13,000,000 26.0 12,000,000 24.0 13,000,000 26.0  Including Miscellaneous Equipment (Additional  Transformer, Switchgear, wastewater treatment)  TOTAL PROCESS CAPITAL (TPC) 93,500,000 187 85,880,000 172 85,350,000 17 General Facilities (5% of TPC) 4,675,000 9.4 4,294,000 8.6 4,268,000 8.5  Engineering and Construction Management 9,350,000 18.7 8,588,000 17.2 9,350,000 18.7  Project Contingency (15%) 16,129,000 32.3 14,814,000 29.6 14,845,000 29.7  TOTAL PLANT COST (TPLC) 123,654,000 247.3 113,576,000 227.2 113,813,000 227.6  Allowance for Funds (AFUDC - 7.0% of TPLC) 8,656,000 17.3 7,950,000 15.9 7,967,000 15.9  Owner's Cost (5% of TPLC) 6,183,000 12.0 5,679,000 11.0 5,691,000 11.0  TOTAL PLANT INVESTMENT (TPI) 138,493,000 276.6 127,205,000 254.1 127,471,000 254.6  Inventory Capital (Spare, same as new) 1,004,000 2.0 906,000 1.8 907,000 1.8  Initial Chemicals and Commissioning (same as new) 4.0 1,812,000 2,008,000 3.6 1,814,000 3.6 

 TOTAL CAPITAL REQUIREMENT (TCR) 141,505,000 283 129,923,000 259 130,192,000 260 

Notes:1.0 Accuracy of Estimate +/-20%2.0 Labor cost based on normal 5 x 8 shift operation3.0 Booster ID fan/motor and electrical cost is included. additional tranformers and switchgears are also included4.0 Medium Retrofit Difficulty assumed5.0 Costs are based on year 2006 (No escalation)6.0 Project implementation based on multiple lump sum speciality contracting

Retrofit capital/CAPCOST-wet FGD-2006.xls.xls
National Lime Association 11311-001

Page 1 of 8

Exhibit 5-3: Fixed and Variable O&M Cost/Levelized Costs (500 MW New Unit)

LSFO FGD System

Input for O&M Costs

Appalachian High S Appalachian Low S

1 Number of Operators (40 hrs/wk) 8 8

2 Operating labor Cost, $/hr 50 50

3 Reagent Purity, % 95 95

4 Reagent Stoichiometry 1.03 1.03

5 Reagent Cost, $/ton 15 15

6 Reagent Requirement, t/h 19.59 8.30

7 SO2 Removal Efficiency, % 98 98

8 SO2 Removed, t/h 11.56 4.90

9 By-product Generated - dry, t/h 32.63 13.83

10 By-product Sale Cost - dry, $/ton 0 0

11 Water Requirement, gpm 660 660

12 Water Cost, $/1000 gal 0.75 0.75

13 Aux. Power Requirement, MW 10.5 7.0

14 Aux. Power Cost, $/MWH 45 45

15 Load Factor, % 80 80

16 Waste from Waste water Treatment, t/h 0.4 0.3

17 Waste Disposal cost, $/ton 20.0 20.0

Appalachian High S Appalachian Low S

Fixed O&M Costs

1. Operating Labor Cost ($/yr) $832,000 $832,000

2. Maintenance Materials Cost ($/yr) $1,423,000 $1,284,000

3. Maintenance Labor Cost ($/yr) $949,000 $856,000

4. Administrative and Support Labor ($/yr) $534,000 $506,000

 Total Yearly Fixed O&M Cost = $3,738,000 $3,478,000

Variable Operating Costs

1. Reagent Costs ($/yr) $2,059,000 $873,000

2. Waste Disposal Cost for FGD System ($/yr) $49,000 $38,000

(Dry basis)

3. Credit for Byproduct ($/yr) $0 $0

4. Water Cost ($/yr) $208,000 $208,000

5. Auxiliary Power Costs ($/yr) $3,311,000 $2,208,000

 Total Yearly Variable O&M Cost $5,627,000 $3,327,000

TOTAL YEARLY FIXED AND VARIABLE O&M COSTS $9,365,000 $6,805,000
National Lime Association 11311-001

Page 2 of 8

Exhibit 5-7: Levelized Costs - LSFO System (500 MW New Unit)

Inputs for Levelized Costs

Appalachian High S Appalachian Low S

1 FGD System Life, years 30 30

2 Capital Cost Levelization Factor 14.8 14.8

3 Discount rate, %/yr 9.0 9.0

4 Inflation Rate, % 3.0 3.0

5 Operating Cost Levelization Factor 1.365 1.365

Appalachian High S Appalachian Low S

Total Capital Cost, MM$ 103.4 93.3

Total Fixed O&M Cost, $/yr 3,738,000 3,478,000

Total Variable O&M Cost, $/yr 5,627,000 3,327,000

Levelized capital Cost, MM$/yr 15.30 13.81

Levelized O&M Cost, MM$/yr 12.78 9.29

Total Levelized Cost, MM$/yr 28.09 23.10

Total cents/kW-hr 0.80 0.66
National Lime Association 11311-001

Page 3 of 8

Exhibit 5-4: Fixed and Variable O&M Cost/Levelized Costs (500 MW Retrofit Unit)

LSFO FGD System

Input for O&M Costs

Appalachian High S Appalachian Low S

1 Number of Operators (40 hrs/wk) 12 12

2 Operating labor Cost, $/hr 50 50

3 Reagent Purity, % 95 95

4 Reagent Stoichiometry 1.03 1.03

5 Reagent Cost, $/ton 15 15

6 Reagent Requirement, t/h 19.59 8.30

7 SO2 Removal Efficiency, % 98 98

8 SO2 Removed, t/h 11.56 4.90

9 By-product Generated - dry, t/h 32.63 13.83

10 By-product Sale Cost - dry, $/ton 0 0

11 Water Requirement, gpm 660 660

12 Water Cost, $/1000 gal 0.75 0.75

13 Aux. Power Requirement, MW 10.5 7.0

14 Aux. Power Cost, $/MWH 45 45

15 Load Factor, % 80 80

16 Waste from Waste water Treatment, t/h 0.4 0.3

17 Waste Disposal cost, $/ton 20.0 20.0

Appalachian High S Appalachian Low S

Fixed O&M Costs

1. Operating Labor Cost ($/yr) $1,248,000 $1,248,000

2. Maintenance Materials Cost ($/yr) $1,423,000 $1,284,000

3. Maintenance Labor Cost ($/yr) $949,000 $856,000

4. Administrative and Support Labor ($/yr) $659,000 $631,000

 Total Yearly Fixed O&M Cost = $4,279,000 $4,019,000

Variable Operating Costs

1. Reagent Costs ($/yr) $2,059,000 $873,000

2. Waste Disposal Cost for FGD System ($/yr) $49,000 $38,000

(Dry basis)

3. Credit for Byproduct ($/yr) $0 $0

4. Water Cost ($/yr) $208,000 $208,000

5. Auxiliary Power Costs ($/yr) $3,311,000 $2,208,000

 Total Yearly Variable O&M Cost $5,627,000 $3,327,000

TOTAL YEARLY FIXED AND VARIABLE O&M COSTS $9,906,000 $7,346,000
National Lime Association 11311-001

Page 4 of 8

Exhibit 5-8: Levelized Costs - LSFO System (500 MW Retrofit Unit)

Inputs for Levelized Costs

Appalachian High S Appalachian Low S

1 FGD System Life, years 20 20

2 Capital Cost Levelization Factor 15.7 15.7

3 Discount rate, %/yr 9.0 9.0

4 Inflation Rate, % 3.0 3.0

5 Operating Cost Levelization Factor 1.274 1.274

Appalachian High S Appalachian Low S

Total Capital Cost, MM$ 141.5 129.9

Total Fixed O&M Cost, $/yr 4,279,000 4,019,000

Total Variable O&M Cost, $/yr 5,627,000 3,327,000

Levelized capital Cost, MM$/yr 22.22 20.40

Levelized O&M Cost, MM$/yr 12.62 9.36

Total Levelized Cost, MM$/yr 34.84 29.76

Total cents/kW-hr 0.99 0.85
National Lime Association 11311-001

Page 5 of 8

Exhibit 5-5: Fixed and Variable O&M Cost/Levelized Costs (500 MW New Unit)

MEL FGD System

Input for O&M Costs

Appalachian High S Appalachian Low S

1 Number of Operators (40 hrs/wk) 8 8

2 Operating labor Cost, $/hr 50 50

3 Reagent Purity, % 94 94

4 Reagent Stoichiometry 1.02 1.02

5 Reagent Cost, $/ton 65 65

6 Reagent Requirement, t/h 10.98 4.65

7 SO2 Removal Efficiency, % 98 98

8 SO2 Removed, t/h 11.56 4.90

9 By-product Generated - dry, t/h 32.63 13.83

10 By-product Sale Cost - dry, $/ton 0 0

11 Water Requirement, gpm 660 660

12 Water Cost, $/1000 gal 0.75 0.75

13 Aux. Power Requirement, MW 6.5 5.0

14 Aux. Power Cost, $/MWH 45 45

15 Load Factor, % 80 80

16 Waste from Waste water Treatment, t/h 0.8 0.5

17 Waste Disposal cost, $/ton 20.0 20.0

Appalachian High S Appalachian Low S

Fixed O&M Costs

1. Operating Labor Cost ($/yr) $832,000 $832,000

2. Maintenance Materials Cost ($/yr) $1,272,000 $1,171,000

3. Maintenance Labor Cost ($/yr) $848,000 $780,000

4. Administrative and Support Labor ($/yr) $504,000 $484,000

 Total Yearly Fixed O&M Cost = $3,456,000 $3,267,000

Variable Operating Costs

1. Reagent Costs ($/yr) $5,001,000 $2,119,000

2. Waste Disposal Cost for FGD System ($/yr) $111,000 $64,000

(Dry basis)

3. Credit for Byproduct ($/yr) $0 $0

4. Water Cost ($/yr) $208,000 $208,000

5. Auxiliary Power Costs ($/yr) $2,050,000 $1,577,000

 Total Yearly Variable O&M Cost $7,370,000 $3,968,000

TOTAL YEARLY FIXED AND VARIABLE O&M COSTS $10,826,000 $7,235,000
National Lime Association 11311-001

Page 6 of 8

Exhibit 5-9: Levelized Costs - MEL System (500 MW New Unit)

Inputs for Levelized Costs

Appalachian High S Appalachian Low S

1 FGD System Life, years 30 30

2 Capital Cost Levelization Factor 14.8 14.8

3 Discount rate, %/yr 9.0 9.0

4 Inflation Rate, % 3.0 3.0

5 Operating Cost Levelization Factor 1.365 1.365

Appalachian High S Appalachian Low S

Total Capital Cost, MM$ 93.4 85.8

Total Fixed O&M Cost, $/yr 3,456,000 3,267,000

Total Variable O&M Cost, $/yr 7,370,000 3,968,000

Levelized capital Cost, MM$/yr 13.82 12.69

Levelized O&M Cost, MM$/yr 14.78 9.88

Total Levelized Cost, MM$/yr 28.60 22.57

Total cents/kW-hr 0.82 0.64
National Lime Association 11311-001

Page 7 of 8

Exhibit 5-6: Fixed and Variable O&M Cost/Levelized Costs (500 MW Retrofit Unit)

MEL FGD System

Input for O&M Costs

Appalachian High S Appalachian Low S

1 Number of Operators (40 hrs/wk) 12 12

2 Operating labor Cost, $/hr 50 50

3 Reagent Purity, % 94 94

4 Reagent Stoichiometry 1.02 1.02

5 Reagent Cost, $/ton 65 65

6 Reagent Requirement, t/h 10.98 4.65

7 SO2 Removal Efficiency, % 98 98

8 SO2 Removed, t/h 11.56 4.90

9 By-product Generated - dry, t/h 32.63 13.83

10 By-product Sale Cost - dry, $/ton 0 0

11 Water Requirement, gpm 660 660

12 Water Cost, $/1000 gal 0.75 0.75

13 Aux. Power Requirement, MW 6.5 5.0

14 Aux. Power Cost, $/MWH 45 45

15 Load Factor, % 80 80

16 Waste from Waste water Treatment, t/h 0.8 0.5

17 Waste Disposal cost, $/ton 20.0 20.0

Appalachian High S Appalachian Low S

Fixed O&M Costs

1. Operating Labor Cost ($/yr) $1,248,000 $1,248,000

2. Maintenance Materials Cost ($/yr) $1,272,000 $1,171,000

3. Maintenance Labor Cost ($/yr) $848,000 $780,000

4. Administrative and Support Labor ($/yr) $629,000 $608,000

 Total Yearly Fixed O&M Cost = $3,997,000 $3,807,000

Variable Operating Costs

1. Reagent Costs ($/yr) $5,001,000 $2,119,000

2. Waste Disposal Cost for FGD System ($/yr) $111,000 $64,000

(Dry basis)

3. Credit for Byproduct ($/yr) $0 $0

4. Water Cost ($/yr) $208,000 $208,000

5. Auxiliary Power Costs ($/yr) $2,050,000 $1,577,000

 Total Yearly Variable O&M Cost $7,370,000 $3,968,000

TOTAL YEARLY FIXED AND VARIABLE O&M COSTS $11,367,000 $7,775,000
National Lime Association 11311-001

Page 8 of 8

Exhibit 5-10: Levelized Costs - MEL System (500 MW Retrofit Unit)

Inputs for Levelized Costs

Appalachian High S Appalachian Low S

1 FGD System Life, years 20 20

2 Capital Cost Levelization Factor 15.7 15.7

3 Discount rate, %/yr 9.0 9.0

4 Inflation Rate, % 3.0 3.0

5 Operating Cost Levelization Factor 1.274 1.274

Appalachian High S Appalachian Low S

Total Capital Cost, MM$ 130.2 120.8

Total Fixed O&M Cost, $/yr 3,997,000 3,807,000

Total Variable O&M Cost, $/yr 7,370,000 3,968,000

Levelized capital Cost, MM$/yr 20.44 18.96

Levelized O&M Cost, MM$/yr 14.48 9.91

Total Levelized Cost, MM$/yr 34.92 28.87

Total cents/kW-hr 1.00 0.82