WET FLUE GAS DESULFURIZATION TECHNOLOGY EVALUATION

PROJECT NUMBER 11311-000

PREPARED FOR

NATIONAL LIME ASSOCIATION

JANUARY 2003

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

1.1 Process Chemistry.........................................................................................................................................................3

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........................................................................................................................................................11

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

2. WET FGD PROCESS ADVANTAGES AND DISADVANTAGES COMPARED TO DRY

FGD TECHNOLOGY............................................................................................................................ 13

2.1 Process Advantages.....................................................................................................................................................13

2.2 Process Disadvantages................................................................................................................................................13

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

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

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

3.2.1 Reagent Handling and Preparation Systems.........................................................................................................17
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3.2.2 SO2 Removal System ...........................................................................................................................................18

3.2.3 Flue Gas System/Stack.........................................................................................................................................19

3.2.4 By-Product Treatment and Handling System.......................................................................................................20

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 SO2 Reduction .............................................................................................................................................................23

4.5 Reagent Utilization .....................................................................................................................................................24

4.6 By-Product Quality.....................................................................................................................................................24

4.7 Energy Consumption..................................................................................................................................................24

4.8 Retrofit Versus New Units..........................................................................................................................................25

5. COSTS ANALYSES........................................................................................................................ 26

5.1 Capital Costs ...............................................................................................................................................................26

5.2 Operations and Maintenance Costs ..........................................................................................................................27

5.2.1 Fixed O&M Costs................................................................................................................................................27

5.2.2 Variable O&M Costs............................................................................................................................................28

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/Levelized Costs for New Units – LSFO FGD System

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

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

5-6 Fixed and Variable O&M Cost/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: January 17, 2003

Raj Gaikwad Date

Technical Advisor

Reviewed by: Willard L. Boward January 17, 2003

Technical Advisor Date

Approved by: January 17, 2003

William DePriest Date

Project Director
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1. FLUE GAS DESULFURIZATION (FGD) DESCRIPTION

Wet FGD technology, which is based on using limestone or lime as a reagent, is a wet scrubbing process and 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 flue 

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

type of FGD system removes SO2 by scrubbing the flue gas with either a limestone or lime (reagent) slurry. The 

wet FGD process is considered a commercially mature technology and is 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 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 to atomize it to fine droplets and uniformly

contact the gas. The droplets absorb SO2 from the gas, facilitating the reaction of the SO2 with reagent in the

slurry. Hydrogen chloride present in the flue gas is also absorbed and neutralized with reagent, causing an

accumulation of chloride ions in the process liquid. Some of the water in the spray droplets evaporates, cooling 

the gas at the inlet from approximately 300°F to 125°F-130°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 sent to

the stack.

In most wet FGD systems, SO2 collection efficiency is controlled by selecting appropriate design features for the 

system. For example, the quantity of liquid sprayed relative to flue gas is related to the SO2 collection efficiency 

needed and is referred to as 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 type of reagent. An important goal of this 
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study will be to quantify the impact that selection of 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 limestonebased "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 corrosive chloride salts 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 are 

recycled to the process with a portion removed as wastewater for chloride control.
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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 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 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 tendency for gypsum to precipitate on surfaces in the absorber and 

cause plugging of pipes and nozzles by maintaining gypsum concentration in 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 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 $460 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 use of organic acids. Provision is sometimes 

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

organic acids include additional operating cost, uncertainty in long-term supply and pricing of DBA byproducts, 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 increase SO2 capture and allow reduction in power consumption and

equipment costs. When magnesium sulfite is present in slurry in contact with flue gas, it buffers (prevents

from falling sharply) pH of the slurry as it absorbs 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, 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 processes a greater capacity to absorb SO2 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

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

is collected in a scoop from recirculating slurry that has just contacted the flue gas (Figure 2). Slurry

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 [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. 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.
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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.

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 is also used for manufacturing wallboard, as an

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

be directed to a wastewater treatment plant. Magnesium sulfate has no biological oxygen demand, so no

additional (biological) wastewater treatment is required. In this study, wastewater treatment requirements are

assumed to be the same for the MEL and the LSFO processes. No additional cost is assumed for disposing
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magnesium sulfate. 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 recovery reduces salinity and toxic metals content of wastewater, which reduces wastewater

treatment requirements. Magnesium hydroxide has a number of commercial uses, including neutralization of

sulfur trioxide in flue gases in the furnace or at the economizer outlet, thereby reducing corrosion in the air

heater, and lowering sulfuric acid plume potential. 

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. For the LSFO absorber, the higher removal efficiencies are typically attained

through the use of additives such as DBA in the recirculating slurry or by recirculating larger quantities of

slurry (larger L/G ratio) in the absorber. This portion of the study focuses on the use of L/G ratios to achieve 

design SO2 capture rather than being dependent on DBA , which avoids potential by-product contamination

and additional wastewater treatment. 

Most of the scrubbers 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 major FGD process developers, including Alstom, Mitsubishi, Babcock & Wilcox, and

Wheelabrator, have shown that a 130 L/G is adequate to achieve 98% efficiency in a typical open-spray 

tower design on the 4.72 lb/MBtu inlet SO2 basis (high-sulfur coal case) of this work. This is further verified 

by recent guarantees offered by FGD vendors for new unit applications. It should be noted that 98% SO2

removal for a LSFO system is, in most cases, not the optimal design point for the process. Efficiencies above 

95% require exponentially greater mass transfer (e.g., L/G) capability and will only be pursued if air permit

requirements necessitate very low SO2 emissions. Note also that these very low SO2 emissions have been

required on some recent air permit applications.
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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 LSFO absorber design. Limestone reagent utilization of at least 97% is necessitated by gypsum

quality requirements. 

Regarding emission averaging period, typically, system design and guarantee requirements are requested to

be demonstrated over a short and manageable test period. This test period is done usually on a one-time or

two-time basis, and consists of 3 two- to four-hour tests. Margin is specified into the design (efficiency,

equipment redundancy, etc.) to meet the plant’s emission averaging 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 

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 fuels, 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. Applications include upgrading commercial designs 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:

• Alstom

• Babcock & Wilcox

• Babcock Power, Inc.
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• Chiyoda

• Hamon Research Cottrell

• Mitsubishi

• Wheelabrator Air Pollution Control
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2. WET FGD PROCESS ADVANTAGES AND DISADVANTAGES COMPARED TO 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 98% can be attained.

3. Adequate and commercially viable suppliers offer the technology.

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

5. Waste gypsum is stable for landfills 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 sufficiently sized gypsum market near the plant.

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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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

High 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

* All sulfur in coal is assumed to be converted to SO2 (no SO3), and that there is no sulfur removed in the

either the bottom ash or electrostatic precipitator ash.

** The air 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 MWnet 2.0 1.3 1.4 1.0

Reagent Limestone Limestone MEL MEL

Reagent cost, $/ton 15 15 50 50

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 14.5 14.5 14.5 14.5

Retrofit application 15.43 15.43 15.43 15.43

Discount rate, % 8.75 8.75 8.75 8.75

Inflation rate, % 2.5 2.5 2.5 2.5

Operating cost levelization factor:

New unit application 1.30 1.30 1.30 1.30

Retrofit application 1.22 1.22 1.22 1.22

*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 16-hour capacity

day bin. The limestone day bin and a gravimetric feeder supply the limestone to a 150% capacity ball mill

system. This allows two-shift operation of the reagent preparation system with the boiler continuously operating 

at 100% load. 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 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 16-hour capacity day bin. The lime day bin and a gravimetric feeder supply the lime to a 150%

capacity slaking system. This will allow two-shift operation of the reagent preparation system with the boiler 

continuously operating at 100% load. A modern, conventional lime slaker with a high efficiency grit removal 

and lime recovery system is used. Two 100% slurry transfer pumps are used to provide high reliability to

transfer the slurry to the slurry tank. Process makeup water is added to the slaker classification 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.
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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

open-spray 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 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 solid 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 oxidized 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 317 LMN steel. The chloride levels 

will be maintained to less than 8,000 ppm. 

The majority of water loss is through evaporation, which is assumed to be the same for both processes. 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 same water losses in the gypsum by-product. The 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

open-spray 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 
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L/G designs. The absorber is constructed of 317 LMN steel or equivalent. The chloride levels will be

maintained to less than 8,000 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 provided 

to direct a small portion of the recirculated slurry to an oxidation tank for the conversion of CaSO3 to CaSO4. 

The remaining slurry falls into the recycle tank. 

 The oxidation tank is sized for 4 hours of solid 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. 

An Oxallizer, by EIMCO, is currently being demonstrated and has the potential to reduce overall costs. 

EIMCO has installed four Oxallizers at the Lowman power station, which uses the MEL process. Those

Oxallizers produce a total of 12-15 t/hr of gypsum, using a bleed stream from the absorber. An Oxallizer

unit is also being demonstrated at a confidential site. The Oxallizers may effectively compete against sparged 

oxidizers and, thus, must be given consideration to determine the overall capital and O&M costs during the

process selection. 

S&L contacted EIMCO to discuss the application of the Oxallizers process to the hypothetical applications of 

the MEL process in this study. On a total cost of electricity impact basis, the Oxallizers process is

approximately 10% more costly than the conventional process under the assumptions used in this analysis. 

 The capital costs presented in this report are based on the conventional sparged oxidation system rather than the 

Oxallizer.

3.2.3 Flue Gas System/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 acid brick chimney. The draft requirements of the LSFO FGD system

will be accommodated by the booster ID fans. The booster ID fan is sized to provide an additional 10" H2O (9" 
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operating) and 9” H2O (8" operating) pressure drop through the absorber for high- and low-sulfur 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 bleeding will be constructed of a carbon steel

substrate with a 1/16-inch “wallpaper” 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” H2O (7” operating).

3.2.4 By-Product Treatment and Handling System

The by-product system will be similar both for 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% solid, is fed to a belt filter for secondary dewatering. Two 

100% 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 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 stream can be disposed of in an onsite landfill/pond or used as a wetting agent for ash

transportation.

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

covered storage area. The covered storage area is sized for 7 days. The gypsum is subsequently 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
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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.

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 (16 hours’ storage)

Limestone crushers

Limestone ball mills (150% 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 (16 hours’ storage)

Lime slakers (150% 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:
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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

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/dampers

Acid brick lined stack for retrofit applications

Bypass duct to existing stack

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 fuel gas flow from a 1000 MW bituminous coal-fired plant are

commercially demonstrated and available from industry suppliers.

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

The SO2 removal guarantees of up to 98% are available from the system suppliers and have been

demonstrated in commercial applications.

4.4 SO2 REDUCTION

Suppliers of FGD systems have guaranteed SO2 reduction efficiency up to 98% with the inlet SO2

concentration up to 10.0 lb/MBtu. To the best of our knowledge, the lowest outlet SO2 loading guaranteed by 

a supplier of FGD systems is 0.06 lb/MBtu.
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4.5 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

manufacturing. Lime reagent utilization is near 100% due to high reactivity of lime and small particle size of 

slaked lime (3 microns average). 

“Stoichiometry” is a related measure of reagent utilization. Stoichiometry is the ratio of reagent fed to SO2

absorbed (mole/mole). Stoichiometry is equal to 100 divided by reagent utilization. For example, a reagent 

utilization of 97% is equal to a stoichiometry of 1.03. 

4.6 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.

4.7 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 71% lower (highsulfur case) than for the LSFO process and the power requirement for the booster ID fan is 22% lower (highsulfur case) than for the LSFO process. Overall, the MEL process will require approximately 0.6% less

auxiliary (% of MWnet) power for high-sulfur coal and approximately 0.3% less auxiliary power auxiliary (% 

of MWnet) for low-sulfur coal compared to the LSFO process. 
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4.8 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.

Most retrofit units will not be able to use existing stacks, as these stacks are typically designed for hot flue

gas approximately at a 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 stack. 
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5. COSTS ANALYSES

5.1 CAPITAL COSTS

Estimated capital costs for the high- and low-sulfur coal applications both of 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 during construction (AFUDC)

• Initial inventory and spare parts (1% of the process capital)

• 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. 

In many cases, owners will also place capital financing charges in this category. We have assumed in our

analysis that the capital cost estimate is an “overnight” cost and therefore no financing charges or escalation 

are included. 
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Salient features of each capital cost estimate are described in Section 3 of this report.

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 contingency, pass-through mark-up, and 

profit.

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. It 

is assumed the plant layout for retrofit application is not 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-labor-efficient 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. 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 

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

various electrical and control users typically needed for FGD operations.

The FGD system auxiliary power requirement has two types of impact on a coal plant:

1. An increase in the net plant heat rate, resulting in higher annual consumption of fuel and the other

materials and services that are directly proportional to furnace heat input, such as any required FGD

reagents and ash disposal.

2. A reduction in net plant MW capacity.

The economic impacts of Item 1 are a direct function of the delivered fuel price, gross turbine heat rate,

boiler efficiency, coal ash content, coal sulfur content, ash disposal cost, reagent costs, by-product values, 

etc.

The economic impact of Item 2, can be considered either from the standpoint of an independent power

producer (IPP) on a competitive grid or from the standpoint of an integrated electric utility operating a power 

grid under regulation. The economic impact would essentially be identical under an IPP or integrated utility

scenario; however, each situation must be considered on a case-by-case basis.

From the IPP point of view, the capacity reduction represents (a) loss of any capacity revenues (including 

ancillary services associated with reserve margin support) that the plant would have received during the

hours when the plant was operating at full load, plus (b) loss of contribution margin during that period

represented by the hourly marginal price of power on the system, less variable operating costs of the unit

under study. From the integrated utility point of view, the economic losses include the cost of buying

replacement capacity and replacement energy, where the net replacement energy cost would be the market
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price of power, less the variable costs (primarily fuel costs) that would be avoided. Either value will be a

function of the hours per year that the plant operates at 100% load. The hours per year that the coal plant

will operate at full load are a function of the load-duration curve of the system the plant is serving and the

competitiveness of this plant relative to other plants connected to this grid. For this reason, each situation

must be considered individually in the context of those parameters.

In the short term, capacity may have very little value in some markets, as indicated Table 5.2.2 -1 below of

recent total spot market costs for electricity: 

TABLE 5.2.2 - 1

RECENT TOTAL SPOT MARKET ELECTRICITY COSTS

$/MWh

Region On Peak Off Peak

Western United States:

Mid-Columbia 26.42 22.67

CA-OR Border 31.19 24.42

NP 15 38.22 26.26

SP 15 37.05 25.42

Palo Verde 29.24 20.15

Gulf/Southeast:

SPP 29.03 14.57

ERCOT 33.75 24.50

SERC 28.72 16.58

FRCC 31.50 21.00

Midwest:

ECAR 26.72 14.10

MAIN 27.59 13.71

MAPP 31.99 16.26

East Coast:
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NEPOOL 46.75 34.00

NY Zone J 68.38 41.70

PJM West 28.62 18.33

The above table shows a range of slightly less than $15/MWh to just over $40/MWh, with an average of

$20/MWh for the off peak periods (most of the hours of the day). Over the longer term, electricity prices are 

likely to be higher in most regions, and a portion of those prices will be attributable to the value of capacity. 

For an optimized power system, the market value of capacity typically will be approximately equal to the

carrying charges on the lowest cost type of capacity (simple cycle combustion turbines), equal to about

$55/kW/year. Depending on unit size, fuel, financing, etc., the levelized total cost of electricity for a modern 

coal plant will be in the neighborhood of $45/MWh, covering capital, O&M, and fuel costs. Market prices

for base-load electricity must reach that general level over the long term for cost-effective installation of new 

coal-fired generating plants.

Although local power prices are depressed in some markets today, for long-term FGD studies involving new 

coal-fired plants that are expected to be economically viable over the long term, a reasonable assumption for 

the cost of auxiliary power in feasibility studies is the full-cost-recovery figure for the plant being studied

(i.e., $45/MWh). This implies that that cost is adequate for installation of the power plant to be viable in the 

first place. This is the appropriate auxiliary power penalty because those revenues are what would be

forgone due to the auxiliary power impact. The forgone revenues for a very old coal plant backfitted with an 

FGD system could, arguably, be lower, since the required capacity charges would be lower. 

Note that this full-cost methodology is what EPRI included in its Technical Assessment Guide when that

guide was only published in paper format.

Considering these generalized conditions, Sargent & Lundy believes that $30/MWh is fairly representative of 

the future mix of the power market. We would propose using a range of $20/MWh to $45/MWh to

demonstrate the sensitivity of the technology selection to this variable. 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. 
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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.

5.3 LEVELIZED COSTS

Levelized costs, also referred to as “life cycle costs,” take into account the impacts of capital costs and O&M 

cost during the operation of a plant over the period of analysis. The levelized fixed charge rate (impact due 

to capital cost) was calculated based on an assumption that a typical customer 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 an 8.75% discount rate and 30-year or 20-year life 

expectancy for new or retrofit facilities, respectively, the levelized fixed charge rates are 14.50% (30-year 

life) and 15.43% (20-years life). The levelized 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 were 1.30 for the 30-year period and 1.22 for the 20-year analysis.
WET FLUE GAS DESULFURIZATION

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DECEMBER 16, 2002

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Project Number 11311-000

A-1

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.)*

1 1 1 1

Absorber inlet expansion 

(“w.c.)

0.5 0.5 0.5 0.5

Absorber spray (“w.c.) 3.12 1.92 0.96 0.72

L/G 130 80 40 30

“w.c. loss per 25 L/G 

(assumed)

0.6 0.6 0.6 0.6

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.)

2 2 2 2

Margin (“w.c.) 1 1 1 1

Total 9.1 7.9 7.0 6.7

*”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) 530 225 330 140

SO2 Absorption (kW) 3391 2010 990 742

No. operating spray 

levels

4 3 2 2

#4 (highest) header 

elevation (ft)

52.5 -- -- --

#3 header elevation 

(ft)

47.5 47.5 -- --

#2 header elevation 

(ft)

42.5 42.5 42.5 42.5

#1 (lowest) header 

elevation (ft)

37.5 37.5 37.5 37.5

Average header 

elevation (ft)

45 42.5 40 40

Nozzle pressure loss 

(psi)

10 10 10 10

Specific gravity of 

pumped slurry

1.09 1.09 1.16 1.16

GPM slurry 196000 120615 58000 43500

Pump efficiency (%) 85.00 85.00 85.00 85.00

Motor efficiency (%) 95 95 95 95

Oxidation Air Compression 

(kW)

2235 947 2235 947

Flue Gas Handling (kW) 2420 2042 1882 1765

By-product Handling (kW) 420 178 420 178

General (kW) 60 60 60 60

Total (kW) 9056 5462 5917 3832
WET FLUE GAS DESULFURIZATION

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A-3

Parasitic load (% of 

generation)

1.81 1.09 1.18 0.77

Margin (%) 0.2 0.2 0.2 0.2

Total (%) 2.01 1.29 1.38 0.97

Parasitic load used in 

study(%)

2.00 1.30 1.40 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 49,000 gpm show a somewhat higher pump efficiency

of 87-89%, whereas the MEL pump curves at 29,000 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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B-1

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 )

2 2 3

3 3 3 2

- +

- +

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

3 2 2

3 2

3 2

3 2 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 3in solution

Flue gas

with SO2

Cleaned

Gas

To Forced

Oxidation

& Gypsum

Production

Slaked Lime

94 moles

Ca(OH)

6 moles

Mg(OH)

2

2

M

CaSO 3 2 1⁄2 H O

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):
WET FLUE GAS DESULFURIZATION

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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
WET FLUE GAS DESULFURIZATION

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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
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B-4

the absorber via 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 11311-000

1/22/03

Exhibit 5-1: Capital Cost for New Units using Appalachian Coals - Wet FGD

Limestone Forced Oxidation Mangesium Enhanced Lime

Appalachian High Sulfur Appalachian Low Sulfur Appalachian High Sulfur Appalachian Low Sulfur

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

 Reagent Feed System 7,373,000 14.7 6,215,000 12.4 5,600,000 11.2 4,824,000 9.6

 SO2 Removal System 17,712,000 35.4 15,000,000 30.0 14,290,000 28.6 13,011,000 26.0

 Flue Gas System 7,857,000 15.7 7,250,000 14.5 7,769,000 15.5 7,150,000 14.3

 By-product Treatment and Handling 7,419,000 14.8 6,018,000 12.0 7,419,000 14.8 6,018,000 12.0

 General Support Equipment 2,011,000 4.0 1,710,000 3.4 2,011,000 4.0 1,710,000 3.4

 (Inclu. miscellaneous Equipment)

 TOTAL PROCESS CAPITAL (TPC) 42,372,000 85 36,193,000 72 37,089,000 74 32,713,000 65

 General Facilities (5% of TPC) 2,119,000 4.2 1,810,000 3.6 1,854,000 3.7 1,636,000 3.3

 Engineering and Construction Management 4,237,000 8.5 3,619,000 7.2 3,709,000 7.4 3,271,000 6.5

 Project Contingency (15%) 7,309,000 14.6 6,243,000 12.5 6,398,000 12.8 5,643,000 11.3

 TOTAL PLANT COST (TPLC) 56,037,000 112.1 47,865,000 95.7 49,050,000 98.1 43,263,000 86.5

 Allowance for Funds (AFUDC - 3.2% of TPLC) 1,793,000 3.6 1,532,000 3.1 1,570,000 3.1 1,384,000 2.8

 Owner's Cost (5% of TPLC) 2,802,000 6.0 2,393,000 5.0 2,453,000 5.0 2,163,000 4.0

 TOTAL PLANT INVESTMENT (TPI) 60,632,000 121.7 51,790,000 103.8 53,073,000 106.2 46,810,000 93.3

 Inventory Capital (Spare, 1% of TPI) 606,000 1.2 518,000 1.0 531,000 1.1 468,000 0.9

 Initial Chemicals and Commissioning (2% of TPI) 1,213,000 2.4 1,036,000 2.1 1,061,000 2.1 936,000 1.9

 Royalties 0 0 0 0 0 0 0 0

 TOTAL CAPITAL REQUIREMENT (TCR) 62,451,000 125 53,344,000 107 54,665,000 109 48,214,000 96

Notes:

1.0 Accuracy of Estimate +-20%

2.0 Labor cost based on normal 5 x 8 shift operation

3.0 Project implementation based on multiple lump sum specialty contracting

New capital/CAPCOST-wet FGD.xls
National Lime 11311-000

1/22/03

Exhibit 5-2: Capital Cost for Retrofit Units using Appalachian Coals - Wet FGD

Limestone Forced Oxidation Mangesium Enhanced Lime

AppalachianHigh Sulfur Appalachian Low Sulfur AppalachianHigh Sulfur Appalachian Low Sulfur

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

 Reagent Feed System 7,588,000 15.2 6,400,000 12.8 6,114,000 12.2 4,970,000 9.9

 SO2 Removal System 22,133,000 44.3 18,540,000 37.1 17,100,000 34.2 15,679,000 31.4

 Flue Gas System with new stack 15,379,000 30.8 15,370,000 30.7 15,264,000 30.5 15,260,000 30.5

 By-product Treatment and Handling 7,900,000 15.8 6,490,000 13.0 7,900,000 15.8 6,490,000 13.0

 General Support Equipment 5,800,000 11.6 5,300,000 10.6 5,800,000 11.6 5,300,000 10.6

 Including Miscellaneous Equipment (Additional 

 Transformer, Switchgear) 

 TOTAL PROCESS CAPITAL (TPC) 58,800,000 118 52,100,000 104 52,178,000 104 47,699,000 95

 General Facilities (5% of TPC) 2,940,000 5.9 2,605,000 5.2 2,609,000 5.2 2,385,000 4.8

 Engineering and Construction Management 5,880,000 11.8 5,210,000 10.4 5,218,000 10.4 4,770,000 9.5

 Project Contingency (15%) 10,143,000 20.3 8,987,000 18.0 9,001,000 18.0 8,228,000 16.5

 TOTAL PLANT COST (TPLC) 77,763,000 155.5 68,902,000 137.8 69,006,000 138.0 63,082,000 126.2

 Allowance for Funds (AFUDC - 3.2% of TPLC) 2,488,000 5.0 2,204,864 4.4 2,208,000 4.4 2,018,624 4.0

 Owner's Cost (5% of TPLC) 3,888,000 8.0 3,445,000 7.0 3,450,000 7.0 3,154,000 6.0

 TOTAL PLANT INVESTMENT (TPI) 84,139,000 168.5 74,551,864 149.2 74,664,000 149.4 68,254,624 136.2

 Inventory Capital (Spare, same as new) 606,000 1.2 518,000 1.0 531,000 1.1 468,000 0.9

 Initial Chemicals and Commissioning (same as new) 1,213,000 2.4 1,036,000 2.1 1,061,000 2.1 936,000 1.9

 Royalties 0 0 0 0 0 0 0 0

 TOTAL CAPITAL REQUIREMENT (TCR) 85,958,000 172 76,105,864 152 76,256,000 153 69,658,624 139

Notes:

1.0 Accuracy of Estimate +-20%

2.0 Labor cost based on normal 5 x 8 shift operation

3.0 Booster ID fan/motor and electrical cost is included

 additional tranformers and switchgears are also included

4.0 Medium Retrofit Difficulty assumed

5.0 Project implementation based on multiple lump sum specialty contracting

Retrofit capital/CAPCOST-wet FGD.xls
National Lime 11311-000

Page 1 of 1

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

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 10.0 6.5

14 Aux. Power Cost, $/MWH 30 30

15 Load Factor, % 80 80

Appalachian High S Appalachian Low S

Fixed O&M Costs

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

2. Maintenance Materials Cost ($/yr) $1,525,000 $1,303,000

3. Maintenance Labor Cost ($/yr) $1,017,000 $869,000

4. Administrative and Support Labor = $555,000 $510,000

 Total Yearly Fixed O&M Cost = $3,929,000 $3,514,000

Variable Operating Costs

1. Reagent Costs = $2,059,000 $873,000

2. Waste Disposal Cost for FGD System = $0 $0

(Dry basis)

3. Credit for Byproduct = $0 $0

4. Water Cost= $208,000 $208,000

5. Additional Power Costs = $2,102,000 $1,367,000

 Total Yearly Variable O&M Cost = 4,369,000 2,448,000

TOTAL YEARLY FIXED AND VARIABLE O&M COSTS 8,298,000 5,962,000

Levelized Costs 

Inputs for Levelized Costs

Appalachian High S Appalachian Low S

1 FGD System Life, years 30 30

2 Capital Cost Levelization Factor 14.5 14.5

3 Discount rate, %/yr 8.75 8.75

4 Inflation Rate, % 2.5 2.5

5 Operating Cost Levelization Factor 1.30 1.30

Total Capital Cost, M$ 62.5 53.3

Levelized capital Cost, MM$/yr 9.06 7.73

Levelized O&M Cost, MM$/yr 10.79 7.75

Total Levelized Cost, MM$/yr 19.84 15.49

Total cents/kW-hr 0.57 0.44
National Lime 11311-000

Page 1 of 1

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

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 10.0 6.5

14 Aux. Power Cost, $/MWH 30 30

15 Load Factor, % 80 80

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) $1,525,000 $1,303,000

3. Maintenance Labor Cost ($/yr) $1,017,000 $869,000

4. Administrative and Support Labor = $680,000 $635,000

 Total Yearly Fixed O&M Cost = $4,470,000 $4,055,000

Variable Operating Costs

1. Reagent Costs = $2,059,000 $873,000

2. Waste Disposal Cost for FGD System = $0 $0

(Dry basis)

3. Credit for Byproduct = $0 $0

6. Water Cost= $208,000 $208,000

7. Additional Power Costs* = $2,102,000 $1,367,000

 Total Yearly Variable O&M Cost = 4,369,000 2,448,000

TOTAL YEARLY FIXED AND VARIABLE O&M COSTS 8,839,000 6,503,000

Levelized Costs 

Inputs for Levelized Costs

Appalachian High S Appalachian Low S

1 FGD System Life, years 20 20

2 Capital Cost Levelization Factor 15.43 15.43

3 Discount rate, %/yr 8.75 8.75

4 Inflation Rate, % 2.5 2.5

5 Operating Cost Levelization Factor 1.22 1.22

Total Capital Cost, M$ 86.0 76.1

Levelized capital Cost, MM$/yr 13.26 11.74

Levelized O&M Cost, MM$/yr 10.78 7.93

Total Levelized Cost, MM$/yr 24.05 19.68

Total cents/kW-hr 0.69 0.56
National Lime 11311-000

Page 1 of 1

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

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 50 50

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 7.0 5.0

14 Aux. Power Cost, $/MWH 30 30

15 Load Factor, % 80 80

Appalachian High S Appalachian Low S

Fixed O&M Costs

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

2. Maintenance Materials Cost ($/yr) $1,335,000 $1,178,000

3. Maintenance Labor Cost ($/yr) $890,000 $785,000

4. Administrative and Support Labor = $517,000 $485,000

 Total Yearly Fixed O&M Cost = $3,574,000 $3,280,000

Variable Operating Costs

1. Reagent Costs = $3,847,000 $1,630,000

2. Waste Disposal Cost for FGD System = $0 $0

(Dry basis)

3. Credit for Byproduct = $0 $0

6. Water Cost= $208,000 $208,000

7. Additional Power Costs* = $1,472,000 $1,051,000

 Total Yearly Variable O&M Cost = 5,527,000 2,889,000

TOTAL YEARLY FIXED AND VARIABLE O&M COSTS 9,101,000 6,169,000

Levelized Costs 

Inputs for Levelized Costs

Appalachian High S Appalachian Low S

1 FGD System Life, years 30 30

2 Capital Cost Levelization Factor 14.5 14.5

3 Discount rate, %/yr 8.75 8.75

4 Inflation Rate, % 2.5 2.5

5 Operating Cost Levelization Factor 1.3 1.3

Total Capital Cost, M$ 54.7 48.2

Levelized capital Cost, MM$/yr 7.93 6.99

Levelized O&M Cost, MM$/yr 11.83 8.02

Total Levelized Cost, MM$/yr 19.76 15.01

Total cents/kW-hr 0.56 0.43
National Lime 11311-000

Page 1 of 1

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

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 50 50

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 7.0 5.0

14 Aux. Power Cost, $/MWH 30 30

15 Load Factor, % 80 80

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) $1,335,000 $1,178,000

3. Maintenance Labor Cost ($/yr) $890,000 $785,000

4. Administrative and Support Labor = $641,000 $610,000

 Total Yearly Fixed O&M Cost = $4,114,000 $3,821,000

Variable Operating Costs

1. Reagent Costs = $3,847,000 $1,630,000

2. Waste Disposal Cost for FGD System = $0 $0

(Dry basis)

3. Credit for Byproduct = $0 $0

6. Water Cost= $208,000 $208,000

7. Additional Power Costs* = $1,472,000 $1,051,000

 Total Yearly Variable O&M Cost = 5,527,000 2,889,000

TOTAL YEARLY FIXED AND VARIABLE O&M COSTS 9,641,000 6,710,000

Levelized Costs 

Inputs for Levelized Costs

Appalachian High S Appalachian Low S

1 FGD System Life, years 20 20

2 Capital Cost Levelization Factor 15.43 15.43

3 Discount rate, %/yr 8.75 8.75

4 Inflation Rate, % 2.5 2.5

5 Operating Cost Levelization Factor 1.22 1.22

Total Capital Cost, M$ 76.3 69.7

Levelized capital Cost, MM$/yr 11.77 10.75

Levelized O&M Cost, MM$/yr 11.76 8.19

Total Levelized Cost, MM$/yr 23.53 18.93

Total cents/kW-hr 0.67 0.54