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P REPARED FOR 

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P REPARED BY

55 East Monroe Street Chicago, IL 60603-5780 USALEGAL 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 th e use of any information or methods disclosed in this report or (b) assumes any liability with resp ect to the use of any information or methods disclosed in this report.

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### C ONTENTS 

### S ECTION 

### P AGE

Wet FGD-2006.doc Project Number 11311-001 i 

### 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 SO 2 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 Desi gn Criteria.................................................................................................... ..........................................14 3.2 System Components ........................................................................................................... ........................................16 3.2.1 Reagent Handling and Preparation Systems .................................................................................. ..........................17 3.2.2 SO 2 Removal System ................................................................................................................ ...............................17 3.2.3 Flue Gas Sy stem and Stack ................................................................................................. .....................................19 3.2.4 Gypsum Dewatering and Handling System ..................................................................................... ........................20

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Wet FGD-2006.doc Project Number 11311-001 ii 3.2.5 General Plan t Support Systems ............................................................................................. ...................................20



### 4. TECHNOLOGY APPLICAT ION CONSTRAINTS ................................................................... 23

4.1 Unit/Abso rber 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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Wet FGD-2006.doc Project Number 11311-001 iii E XHIBITS N UMBER T ITLE 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

A PPENDIX N UMBER T ITLE A Auxiliary Power Summary B Detailed Description of MEL Process Chemistry

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Wet FGD-2006.doc Project Number 11311-001 iv R EPORT P REPARED , R EVIEWED , AND A PPROVED BY S ARGENT & L UNDY 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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Wet FGD-2006.doc Project Number 11311-001 1 

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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 (SO 2 ) reduction from coal-fired utility boilers. The wet FGD flue gas treatment system is typi cally located after removal of particulate matter from the flue gas either by a baghouse or by an electrostatic precip itator. The cleaned gas is di scharged 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 fluegas stream through a limestone or lime slurry spray. In typical absorber designs, the gas flows upward through th e absorber countercurrent to the spray liquor flowing downward through the absorber. However, other desi gns are also available, including co-current and countercurrent designs, and designs where the gas is fo rced 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 SO 2from the gas, facilitating the reaction of the SO 2with the reagent in the slurry. Hydrogen chlorideand 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 dropl ets 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 SO 2collection efficiency required is achieved by selecting a combination of design features for the system. For example, one featur e is the quantity of liquid sprayed relative to flue gas flow and is referred to as the liquid-to-g as (L/G) ratio. Higher L/G ratios improve SO 2removal by exposing the gas to more absorbing liquor. However, higher L/G ratio s 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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Wet FGD-2006.doc Project Number 11311-001 2 absorber pressure drop and quantity/type of reagent. An im portant 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 th e process is replenished by adding fresh limestone or lime slurry to the reaction tank. In the LSFO process, the slurry is al so aerated in the reaction tank to oxidize calcium sulfite hemihydrate (CaSO 3 • 1⁄2 H 2 O) to calcium sulfate dihydrate (CaSO 4 • 2H 2 O), 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 thenrecirculated 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 produ ce 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 a nd 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 re moved 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 sim ilar to an limestone forced oxidation system. The MEL system is typically considered to be the more attractiv e wet lime system of the two options and is therefore the process evaluated in this study.

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Wet FGD-2006.doc Project Number 11311-001 3 1.1 PROCESS CHEMISTRY
1.1.1 LSFO Process A simplified explanation of the LSFO process (Figure 1) is that the SO 2absorbed in the reci rculated slurry reacts with dissolved limestone (CaCO 3 ) in the slurry to form calcium sulfite hemihydrate (CaSO 3 • 1⁄2H 2 O) according to the following simplified reaction:

SO 2+ CaCO 3 + 1⁄2 H 2 O ⇒→ CaSO 3 • 1⁄2H 2 O + CO 2
Carbon dioxide formed from reaction of limestone with SO 2is released into the flue gas. Oxidation air is bubbled through the slurry to convert CaSO 3 • 1⁄2H 2 O to gypsum (CaSO 4 • 2H 2 O) according to the following simplified reaction:

CaSO 3 • 1⁄2H 2 O + 1⁄2O 2+ 1.5 H 2 O ⇒→ CaSO 4 • 2H 2 O
Aeration oxidizes all the calcium sulfite to calcium sulf ate and forces precipitation to occur on existing gypsum crystals in the reaction tank. This minimizes the tendenc y for gypsum to precipitate on surfaces in the absorber, piping, and nozzles by maintaining the gypsum concentra tion 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 CaSO 3 • 1⁄2H 2 O 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.
SO 2collection 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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Wet FGD-2006.doc Project Number 11311-001 4 Figure 1: LSFO Process SO 2 Absorber



### Recycle Tank Spray Headers

Mist Eliminators

Compressed Air Limestone Slurry Gypsum to Dewatering Gas Inlet Gas Outlet to Stack



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Wet FGD-2006.doc Project Number 11311-001 5 absorbs the acidic SO 2 . This improves absorber performance by increasing solubility of SO 2in 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. (R eagent utilization is the percentage of reagent fed to the FGD system that reacts with SO 2 . 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 absorb er 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 effectiv e. These acids include adipic, glycolic, maleic, acrylic, and formic acids. Dibasic acid (DBA), a by-product fro m 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 dependi ng on the use of organic acids. Provision is sometimes included in designs to allow future addition of organic acid in case SO 2collection efficiency is inadequate. Disadvantages of organic acids include additional operati ng 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 SO 2 (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, SO 2 , which is a strong acid, reacts with and is neutralized by slaked lime. The overall reactions between lime and SO 2 in the absorber are shown in reactions (1) and (2). Calcium hydroxide in the slurry reacts with most of the SO 2 to precipitate calcium sulfite (CaSO 3 • 1⁄2 H 2 O). Magnesium hydroxide reacts with the remainder of the SO 2 to form soluble magnesium salts; magnesium sulfite and magnesium bisulfite [MgSO 3 , Mg(HSO 3 ) 2 ].

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Wet FGD-2006.doc Project Number 11311-001 6 Figure 2: MEL Process SO 2 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
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Wet FGD-2006.doc Project Number 11311-001 7 These soluble magnesium salts greatly enhance SO 2 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 SO 2 . This improves absorber performance by increasing solubility of SO 2in the slurry, which can allow operation at a lowe r 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 SO 2 per specific quantity of absorbing liquor than LSFO. Since magnesium sulfite performs the same function as d escribed for organic acids in LSFO systems, no organic acid addition is required to improve SO 2collection 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 su lfate. 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 a bout 10% of the concentration at which scale would form.

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

(2) 4SO 2 + 3Mg(OH) 2 → 2MgSO 3 + Mg(HSO 3 ) 2 +2H 2 O

A portion of the recirculated slurry containing absorbed SO 2 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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Wet FGD-2006.doc Project Number 11311-001 8 collected in the scoop has lower pH (~5.5), which mak es it more suitable for the subsequent oxidation step. The slurry is pumped to an external forced oxidation ta nk where it is contacted with air. Calcium sulfite is converted to gypsum and magnesium salts are oxidized to MgSO 4 , according to reactions (3) and (4) & (5), respectively. The gypsum precipitates in a crystalline form while MgSO 4 remains in solution. The gypsum in the slurry effluent from the oxidizerundergoes primary and seconda ry dewatering to produce gypsum by-product cake in the same way as in the LSFO process. Liquid containing MgSO 4 is returned to the absorber after dewatering the gypsum. (3) CaSO 3 • 1⁄2H 2 O + 1⁄2O 2 + 3/2 H 2 O → CaSO 4 • 2H 2 O (4) MgSO 3 + 1⁄2O 2 → MgSO 4 (5) Mg(HSO 3 ) 2 + O 2 → MgSO 4 + H 2 SO 4 

### 1.2 R EAGENTS AND 

### B Y 

### -P RODUCTS
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 u tilization and to minimize the amount of unreacted limestone in the gypsum product. Some LSFO processes claim to beable 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 SO 2and 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 u sed for making wallboard, as an additive for concrete, or as an agricultural soil conditioner.

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Wet FGD-2006.doc Project Number 11311-001 9 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). Tota l oxides content is equivalent in SO 2 neutralizing value to about 94% CaO. Magnesium-enhanced lime can be produced fro m 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 SO 2 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% su lfur and 0.1% chlorine, 2.3% MgO is required.) Magnesium oxide which reacts with hydrogen ch loride is not available to react with SO 2 to form magnesium sulfite, which has the beneficial effects on SO 2 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 controlle d to a set point near 6.0. At this pH nearly 100% of the lime reacts with SO 2 . As in the LSFO process, the by-product gypsum, after dewateri ng 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 pr oduct or in the recirculated scrubber slurry are directed to a wastewater treatment plant similar to that descr ibed for the LSFO process. In this study, wastewater treatment requirements are assumed to be the same fo r 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 plan t. Magnesium hydroxide has a number of commercial uses, including injection in the flue gas in the furnace, at th e economizer outlet or at the ESP inlet for neutralization of sulfur trioxide (SO 3 ) in flue gases thereby reducing corrosion in th e air heater, and lowering the potential for a visible plume caused by sulfuric acid condensation.

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Wet FGD-2006.doc Project Number 11311-001 10 

### 1.3 SO 2 

### R EMOVAL 

### P ERFORMANCE
1.3.1 LSFO Process LSFO systems have achieved SO 2 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 star t up in 2008 and beyond. For the LSFO absorber, the higher removal efficiencies are typically attained us ing a combination of higher L/G’s, higher velocity, gas/liquid contactors, or through the use of additives su ch as DBA in the recirculating slurry . This study employs a design that is not dependent on DB A 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, Babc ock 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 SO 2 /MBtu inlet) and an L/G of 80 is appropriate for the low sulfur coal case (2.00 lb SO 2 /MBtu inlet). This is further verified by recent guara ntees 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-sulf ur coals, we have estimated that L/G ratios of 130 and 80, respectively, will be required to achieve 98% SO 2removal 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 qua lity requirements.

1.3.2 MEL Process MEL forced oxidation systems have achieved a better level of performance than the LSFO process, with SO 2 removal efficiencies between 98% to 99% in power plants also firing a variety of high- and low-sulfur coals. Because of the higher SO 2absorption capacity available in the magn esium-enhanced lime system compared with

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Wet FGD-2006.doc Project Number 11311-001 11 the LSFO system, we estimate that L/G ratios of 40 and 30 will be requi red to achieve 98% SO 2removal 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 C OMMERCIAL 

### S TATUS Both the LSFO and the MEL FGD systems are operating su ccessfully at many coal-fired power facilities, ranging in size from less than 100 MW to 1000 MW and for coal su lfur content well above 5%. Many competing designs are available in the marketplace, from a number of well-r espected 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 th eir use of the typical mass transfer concepts for achieving high SO 2removal required of modern FGD systems, recent competition in the U.S. market has shown that the various designs are very similar in life cycle co st (balance of capital and operating costs). Therefore, for purposes of this study, we have chosena conventional open spray tower design (currently offered by at least five suppliers) as the basis for this work.

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Wet FGD-2006.doc Project Number 11311-001 12 

## 2. W ET 

## FGD T ECHNOLOGY 

## C OMPARISON WITH 

## D RY 

## FGD T ECHNOLOGY



### 2.1 P ROCESS 

### A DVANTAGES The LSFO and MEL technologies have the following adva ntages when compared with other FGD technologies:
1. Well-established FGD technology on a variety of world coals with proven reliability. 2. SO 2 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 dispo sal without blending with fly ash and lime. 6. Can be designed to produce wallboard-g rade 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 P ROCESS 

### D ISADVANTAGES The LSFO and MEL technologies can have the followi ng 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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Wet FGD-2006.doc Project Number 11311-001 13 5. In most applications, a waste water treatment facility will be required to remove heavy metals from the blowdown stream.

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Wet FGD-2006.doc Project Number 11311-001 14 

## 3. D ESIGN 

## B ASIS



### 3.1 S PECIFIC 

### D ESIGN 

### C RITERIA Table 3.1 -1 lists the specific design criteria.

T ABLE 3.1-1 S PECIFIC D ESIGN C RITERIA 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 SO 2 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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Wet FGD-2006.doc Project Number 11311-001 15 T ABLE 3.1-1 S PECIFIC D ESIGN C RITERIA LSFO LSFO MEL MEL Flue gas temperature at FGD outlet,
° F 130 130 130 130

SO 2 reduction efficiency, % 98 98 98 98 SO 2 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 SO 2 (no SO 3 ) and there is no sulfur removed in the either the botto m ash or electrostatic precipitator ash.

** The air preheater leakage for the new plant and re trofit plants are assumed to be the same over a long duration.
Table 3.1-2 summarizes the parameters used for the FGD comparison.

T ABLE 3.1-2 P ARAMETERS U SED FOR FGD C OMPARISON 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 SO 2 removal, % 98% 98% 98% 98% SO 2 emission, lb/MBtu 0.094 0.04 0.094 0.04 By product Gypsum Gypsum Gypsum Gypsum Power consumption*, % of MW gross 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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Wet FGD-2006.doc Project Number 11311-001 16 T ABLE 3.1-2 P ARAMETERS U SED FOR FGD C OMPARISON LSFO LSFO MEL MEL SO 2 oxidation stoichiometry (O/SO 2 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 gy psum, a reagent stoichiometry (moles reagent per mole SO 2 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 S YSTEM 

### C OMPONENTS 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 descripti ons present important features of the components and assumptions used in this study:

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Wet FGD-2006.doc Project Number 11311-001 17 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 conveyo r 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 recycl e 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 cap acity 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 su lfuric acid feed pumps is also provided.
3.2.2 SO 2 Removal System
3.2.2.1 LSFO Process One 100% capacity absorber is provided to achieve 98% SO 2removal efficiency. The absorber is a vertical openspray tower, with countercurrent contact between th e flue gas and scrubber slurry. To achieve 98% SO 2reduction with high-sulfur coal, 4 operating pumps will be provide d to achieve a 130 L/G ratio. Three operating pumps will

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Wet FGD-2006.doc Project Number 11311-001 18 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 provi ded 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 CaSO 3 to CaSO 4 , a 3:1 stoichiometric ratio of oxygen to absorbed SO 2(moles O/moles SO 2 ) 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 pro cesses 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 chlori de level, which, again, will be the same for both processes. Therefore, the water losses and cons umption will be the same for both wet processes.
3.2.2.2 MEL Process One 100% capacity absorber is provided to achieve 98% SO 2removal efficiency. The absorber is a vertical openspray tower, with countercurrent contact between th e flue gas and scrubber slurry. To achieve 98% SO 2reduction with high-sulfur coal, 2 operating pumps will be provi ded to achieve a 40 L/G ra tio. 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 provide d, along with a dedicated pump, for both L/G designs. The absorber is constructed of UNS S32205 st eel 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 CaSO 3 to CaSO 4 . The remaining slurry returns to the recycle tank.

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Wet FGD-2006.doc Project Number 11311-001 19 The oxidation tank is sized for 4 hours of solids residen ce time. To achieve greater than 99% oxidation of CaSO 3 to CaSO 4 , a 3:1 stoichiometric ratio of oxygen to removed SO 2is used in the oxidation compressor design. Compressed air is injected below the oxidation tank agitato r 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 Oxallizers TM 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 Oxallizers TM process is approximately 10% more costly than the conventional process consider ing 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 Oxallizer TM

. 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" H 2 O (9.4" operating) and 9.2” H 2 O (8.2" operating) pressure drop through the absorber for high- and low-sulfur coal applications, respectively. The inlet ductwork to th e absorber system, including the booster ID fan, will be fabricated with carbon steel. The outlet ductwork from the ab sorber 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” H 2 O (7.1” operating) and 7.8” H 2 O (6.8” operating) for the high- and lowsulfur cases, respectively.

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Wet FGD-2006.doc Project Number 11311-001 20 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 be lt 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 system s, and fire protection systems. These accommodations are included in the design and cost estimate.

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Wet FGD-2006.doc Project Number 11311-001 21 Table 3.2.5–1 lists the equipment used in each subsystem.
TABLE 3.2.5 -1 E QUIPMENT U SED IN E ACH S UBSYSTEM 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) SO 2 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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Wet FGD-2006.doc Project Number 11311-001 22 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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Wet FGD-2006.doc Project Number 11311-001 23 

## 4. T ECHNOLOGY 

## A PPLICATION 

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


### 4.1 U NIT 

### /A BSORBER 

### S IZE 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. Ther efore, 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 C OAL 

### S ULFUR 

### C ONTENT

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

### E XPECTATIONS
SO 2 removal guarantees of up to 99% are available from th e system suppliers and have been demonstrated in commercial applications.


### 4.4 R EAGENT 

### U TILIZATION Reagent utilization is the percentage of reagen t fed to the FGD system that reacts with SO 2 . Reagent utilization in a wet FGD system is typically very close to the requi red theoretical level when considered on a removed SO 2 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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Wet FGD-2006.doc Project Number 11311-001 24 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 B Y 

### -P RODUCT 

### Q UALITY 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 E NERGY 

### C ONSUMPTION The process energy consumption results primarily fro m the booster ID fan power required to overcome the draft loss across the absorber and the power requireme nt for recirculation pumps. The MEL process has an advantage over the LSFO process because the power requireme nt for slurry recirculation is 74% lower (highsulfur case) than for the LSFO process and the power re quirement 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 MW gross ) power for high-sulfur coal and approxima tely 36.9% less auxiliary power auxiliary (% of MW gross ) for low-sulfur coal compared to the LSFO process.


### 4.7 R ETROFIT 

### V ERSUS 

### N EW 

### U NITS In new unit and retrofit applications, wet FGD technol ogy 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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Wet FGD-2006.doc Project Number 11311-001 25 Most retrofit units will not be able to use existing stack s, as these stacks are typically designed for hot flue gas at approximately 100 ft/sec exit velocity. To accommoda te the saturated flue gas from a wet FGD system, wet stacks are typically designed for a highly corrosion-resi stant material with a gas velocity of between 55 to 70 ft/sec. The lower gas velocity is required to prev ent 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 ha mpered by the existing site constraints which may result in unique and not well optimized equipment a rrangements 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 H G 

### R EMOVAL The 1990 Clean Air Act Amendments required the EPA to study the public health risks, if any, associated with hazardous air pollutant emissions, including merc ury, 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 uncertain ties identified in the mercury reports, EPA issued an information collection request (ICR) requiring all coal-fi red 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 obst acle 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 rev ealed 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 inst alled 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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Wet FGD-2006.doc Project Number 11311-001 26 The ICR data indicated that the vaporous form of me rcury 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 stab le 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 east ern bituminous coals had greater removal efficiencies and also tended to have greater concentrations of chlo rides in their coal analysis. The discovery of this correlation has helped define the problem of contro lling 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 exis ting 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 SubBituminous 5 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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Wet FGD-2006.doc Project Number 11311-001 27 This correlation explains many of the differing removal e fficiencies exhibited at power plants. Plants that fire Powder River Basin (PRB) coals have relatively lo w amounts of chloride in the coal, have greater concentrations of elemental mercury, and they exhi bit 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 evaluati on 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 bette r 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 n early 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 consid erable 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 temperatur es. This was evident when evaluators identified that ionic mercury was not read ily 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 r esult has focused investigators efforts at collection at air heater outlet temperatures. While these two correlations resulted from analysis of th e 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 need ed to fully understand mercury speciation and mercury capture and how coal chemistry is related to each issue.

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Wet FGD-2006.doc Project Number 11311-001 28 The availability of the ICR data resulted in an infusi on 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 demonstra tion in a power plant situation. Thus, there is much still remaining with regard to know ing how technologies will perform in a specific application and the reliability of the systems. In conclusion, the preponderance of the data on Me rcury 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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Wet FGD-2006.doc Project Number 11311-001 29 

## 5. C OSTS 

## A NALYSES



### 5.1 C APITAL 

### C OSTS 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 associat ed with the new and retrofit installations of these technologies. The costs were developed primarily from Sargent & L undy’s cost database, which is continuously updated from ongoing work in the area of FGD. The capital cost estimates provided herein represent th e “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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Wet FGD-2006.doc Project Number 11311-001 30 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 contr acting 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 c ontractor’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 O PERATIONS AND 

### M AINTENANCE 

### C OSTS 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, ma intenance 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 un its and 12 operators for the retrofit application. For retrofit applications, we have assumed that the plan t layout and integration between the existing power generation block and the FGD system is not fully op timized, 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, a nd the labor required to perform the maintenance.

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Wet FGD-2006.doc Project Number 11311-001 31 5.2.2 Variable O&M Costs Variable O&M costs developed for each FGD system incl ude 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 asso ciated with the operation of the fans (incremental ID fan power for new units and booste r 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, recirc ulation 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 variab le O&M costs for new and retrofit applications of the LSFO FGD technology, respectively. Exhibit 5-5 a nd Exhibit 5-6 represent the fixed and variable O&M costs for new and retrofit applications of the MEL F GD technology, respectively. Included with each of these Exhibits are the assumptions used for all the co mponents in the fixed and variable O&M costs.


### 5.3 L EVELIZED 

### C OSTS 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% de bt 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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Wet FGD-2006.doc Project Number 11311-001 32 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, respectiv ely. 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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Wet FGD-2006.doc Project Number 11311-001 A-1 APPENDIX A
A UXILIARY P OWER S UMMARY One of the major components of auxiliary power is a dditional power required by ID fans to overcome system pressure drop caused by flue gas flow through the ab sorber. The system pressure drop was estimated as follows in Table A-1:

T ABLE A-1 E STIMATED FGD S YSTEM P RESSURE D ROPLSFO HighSulfur LSFO LowSulfur MEL HighSulfur MEL 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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Wet FGD-2006.doc Project Number 11311-001 A-2 T ABLE A-2 P OWER R EQUIREMENTS WITH F LUE G AS H ANDLING AND O THER FGD S UBSYSTEMSLSFO HighSulfur LSFO LowSulfur MEL HighSulfur MEL LowSulfur Reagent Preparation (kW) 720 305 450 190 SO 2 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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Wet FGD-2006.doc Project Number 11311-001 A-3 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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Wet FGD-2006.doc Project Number 11311-001 B-1 APPENDIX B

D ETAILED D ESCRIPTION OF MEL P ROCESS C HEMISTRY Steps in Absorption of SO 2 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 SO 2 . Slaked lime is added to the recycle tank of an SO 2 absorber to replenish reagent consumed, as shown in Figure 1. Figure 1: Reactions in MEL SO 2 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:



### SO 2 

### SO ABSORBED BY MgSO 2 3 

### Mg(HSO ) 3 2



### MgSO in solution 3 

### Flue gas with SO 2 

### Cleaned Gas

To Forced Oxidation & Gypsum Production Slaked Lim e 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 SO 2 , for a magnesium-enhanced lime with 6 moles of magnesium hydroxide and 94 moles of calci um hydroxide, is shown in reaction (1):

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Wet FGD-2006.doc Project Number 11311-001 B-2 (1) 102 SO 2 + 94 Ca(OH) 2 + 6 Mg(OH) 2 → 94 CaSO 3 • 1⁄2 H 2 O (solid) + 4 MgSO 3 (aqueous) + 2 Mg(HSO 3 ) 2 (aqueous) + 51 H 2 O

Although reaction (1) shows the overall reaction, SO 2 does not react directly with calcium or magnesium hydroxide. SO 2 capture with lime occurs in tw o steps. In the first step, SO 2 is absorbed into the liquid phase of slurry recirculated from the recycle tank thr ough banks of spray nozzles. In the second step, SO 2 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, SO 2 is absorbed by recirculated slurry whose liquid phase contains magnesium sulfite (MgSO 3 ). Magnesium sulfite is alkaline with respect to SO 2 , which is highly acidic. When SO 2 is absorbed into the liquid, it combines with water to form sulf urous acid, which dissociates into hydrogen and hydrogen sulfite ions (HSO 3 - ), shown in reaction (2). MgSO 3 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 SO 2 in the liquid by keeping reaction (1) driven far to the right. By preventing accumulation of SO 2 in the liquid, maximum driving force for mass transfer of SO 2 from flue gas to the liquid is maintained. (2) SO 2 + H 2 O → H + + HSO 3 -

(3) H + + HSO 3 - + MgSO 3 → Mg(HSO 3 ) 2
In the second step, slurry containing captured SO 2 in the form of Mg(HSO 3 ) 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(HSO 3 ) 2 to precipitate calcium sulfite hemihydrate (CaSO 3 • 1⁄2H 2 O) and regenerate MgSO 3 , according to reaction (4). Magnesium hydroxide in the slaked lime reacts with a smaller portion of

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Wet FGD-2006.doc Project Number 11311-001 B-3 Mg(HSO 3 ) 2 to form additional MgSO 3 as shown in reaction (5). A portion of captured SO 2 remains in the form of Mg(HSO 3 ) 2 . (4) Mg(HSO 3 ) 2 + Ca(OH) 2 → CaSO 3 • 1⁄2H 2 O (solid) + MgSO 3 (aqueous) + 11⁄2H 2 O (5) 2Mg(HSO 3 ) 2 + Mg(OH) 2 → 2MgSO 3 (aqueous) + Mg(HSO 3 ) 2 (aqueous) + 2H 2 O A portion of the MgSO 3 formed in reactions (4) and (5) is converted to MgSO 4 by oxygen present in the flue gas. MgSO 4 does not buffer pH in a range effective for SO 2 removal. However, a mechanism called co-precipitation allows a portion of this MgSO 4 to be converted back to MgSO 3 . When CaSO 3 • 1⁄2H 2 O precipitates, it tends to include sulfate (SO 4 = ) ions in its crystal lattice in place of some sulfite (SO 3 = ) ions. (This co-precipitation occurs even when the scrubber liquor is unsaturated with gypsum). As a result, magnesium ions which were associated with the SO 4 = in MgSO 4 become available to pair with SO 3 = and regenerate additional MgSO 3 . Rewriting equation (4) to include this effect, and assuming that the mole fraction of CaSO 4 in the calcium sulfite crystal lattice is 0.2, shows how this additional MgSO 3 is formed: (6) Ca(OH) 2 + Mg(HSO 3 ) 2 + 0.2 MgSO 4 → Ca(SO 3 ) 0.8 (SO 4 ) 0.2 • 1⁄2H 2 O (co-precipitate) + 1.2 MgSO 3 + H 2 O

As a result of this effect, a substantial amount of MgSO 3 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 ju st 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 oxi dation step. Slurry withdrawn from the absorber via

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Wet FGD-2006.doc Project Number 11311-001 B-4 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 SO 2 . The slurry is pumped to an external forced oxida tion tank where it is contacted with compressed air. CaSO 3 • 1⁄2H 2 O is converted to gypsum (CaSO 4 • 2H 2 O), and MgSO 3 is oxidized to MgSO 4 , according to reactions (7) and (8), respectively. The gypsum precipitates while MgSO 4 remains in solution. Liquid containing MgSO 4 is returned to the absorber after dewatering the gypsum. A portion of MgSO 4 returned to the absorber replenishes MgSO 3 as described for reaction (6). Mg(HSO 3 ) 2 in the liquor is oxidized to form MgSO 4 and sulfuric acid (H 2 SO 4 ), shown in reaction (9), which reduces pH to below 5.5 necessary to speed oxidation of CaSO 3 • 1⁄2H 2 O. 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 remain s, which allows overall lime reagent utilization to be more than 99.9%. (7) CaSO 3 • 1⁄2H 2 O + 1⁄2O 2 + 3/2 H 2 O → CaSO 4 • 2H 2 O (gypsum) (8) MgSO 3 + 1⁄2 O 2 → MgSO 4 (9) Mg(HSO 3 ) 2 + O 2 → MgSO 4 + H 2 SO 4 (sulfuric acid) (10) CaO + CaCO 3 +2H 2 SO 4 → 2 CaSO 4 • 2H 2 O + CO 2National Lime Association11311-0015/10/2006Exhibit 5-1: Capital Cost for 500 MW New Unit using Appalachian Coals - Wet FGDLimestone Forced Oxidation Magnesium Enhanced LimeAppalachian High SulfurAppalachian Low SulfurAppalachian High SulfurAppalachian Low SulfurSubsystemsCost, US$ $/kWCost, US$ $/kWCost, US$ $/kWCost, US$ $/kWReagent Feed System 12,825,000 25.7 10,432,800 20.9 8,740,000 17.5 7,866,000 15.7SO2 Removal System 28,475,000 57.0 25,925,000 51.9 25,627,500 51.3 23,332,500 46.7Flue Gas System 11,500,000 23.0 11,270,000 22.5 11,250,000 22.5 11,025,000 22.1By-product Treatment and Handling 9,975,000 20.0 9,025,000 18.1 9,975,000 20.0 9,025,000 18.1General Support Equipment 5,000,000 10.0 4,500,000 9.0 5,000,000 10.0 4,500,000 9.0Including Miscellaneous Equipment (AdditionalTransformer, Switchgear, wastewater treatment)TOTAL PROCESS CAPITAL (TPC) 67,775,000 136 61,152,800 122 60,592,500 121 55,748,500 111General Facilities (5% of TPC) 3,389,000 6.8 3,058,000 6.1 3,030,000 6.1 2,787,000 5.6Engineering and Construction Management 6,778,000 13.6 6,115,000 12.2 6,778,000 13.6 6,115,000 12.2Project Contingency (15%) 11,691,000 23.4 10,549,000 21.1 10,560,000 21.1 9,698,000 19.4TOTAL PLANT COST (TPLC) 89,633,000 179.3 80,874,800 161.7 80,960,500 161.9 74,348,500 148.7Allowance for Funds (AFUDC - 7.0% of TPLC) 6,274,000 12.5 5,661,000 11.3 5,667,000 11.3 5,204,000 10.4Owner's Cost (5% of TPLC) 4,482,000 9.0 4,044,000 8.0 4,048,000 8.0 3,717,000 7.0TOTAL PLANT INVESTMENT (TPI) 100,389,000 200.8 90,579,800 181.1 90,675,500 181.3 83,269,500 166.1Inventory Capital (Spare, 1% of TPI) 1,004,000 2.0 906,000 1.8 907,000 1.8 833,000 1.7Initial Chemicals and Commissioning (2% of TPI)2,008,0004.0 1,812,0003.6 1,814,0003.6 1,665,0003.3TOTAL CAPITAL REQUIREMENT (TCR)103,401,000207 93,297,800 187 93,396,500 187 85,767,500 171TOTAL CAPITAL REQUIREMENT (TCR) withBAGHOUSE128,401,000 257 118,297,800 237 118,396,500 237 110,767,500 222Notes: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 FGDNew capital/CAPCOST-wet FGD-2006.xls.xlsNational Lime Association11311-0015/10/2006Exhibit 5-2: Capital Cost for 500 MW Retrofit Unit using Appalachian Coals - Wet FGDLimestone Forced Oxidation Magnesium Enhanced LimeAppalachian High SulfurAppalachian Low SulfurAppalachian High SulfurAppalachian Low SulfurSubsystemsCost, US$ $/kWCost, US$ $/kWCost, US$ $/kWCost, US$ $/kWReagent Feed System 13,500,000 27.0 11,340,000 22.7 9,200,000 18.4 8,280,000 16.6SO2 Removal System 33,500,000 67.0 30,500,000 61.0 30,150,000 60.3 27,450,000 54.9Flue Gas System with new stack 23,000,000 46.0 22,540,000 45.1 22,500,000 45.0 22,050,000 44.1By-product Treatment and Handling 10,500,000 21.0 9,500,000 19.0 10,500,000 21.0 9,500,000 19.0General Support Equipment 13,000,000 26.0 12,000,000 24.0 13,000,000 26.0 12,000,000 24.0Including Miscellaneous Equipment (AdditionalTransformer, Switchgear, wastewater treatment)TOTAL PROCESS CAPITAL (TPC) 93,500,000 187 85,880,000 172 85,350,000 171 79,280,000 159General Facilities (5% of TPC) 4,675,000 9.4 4,294,000 8.6 4,268,000 8.5 3,964,000 7.9Engineering and Construction Management 9,350,000 18.7 8,588,000 17.2 9,350,000 18.7 8,588,000 17.2Project Contingency (15%) 16,129,000 32.3 14,814,000 29.6 14,845,000 29.7 13,775,000 27.6TOTAL PLANT COST (TPLC) 123,654,000 247.3 113,576,000 227.2 113,813,000 227.6 105,607,000 211.2Allowance for Funds (AFUDC - 7.0% of TPLC) 8,656,000 17.3 7,950,000 15.9 7,967,000 15.9 7,392,000 14.8Owner's Cost (5% of TPLC) 6,183,000 12.0 5,679,000 11.0 5,691,000 11.0 5,280,000 11.0TOTAL PLANT INVESTMENT (TPI) 138,493,000 276.6 127,205,000 254.1 127,471,000 254.6 118,279,000 237.0Inventory Capital (Spare, same as new) 1,004,000 2.0 906,000 1.8 907,000 1.8 833,000 1.7Initial Chemicals and Commissioning (same as new)2,008,0004.0 1,812,0003.6 1,814,0003.6 1,665,0003.3TOTAL CAPITAL REQUIREMENT (TCR)141,505,000283 129,923,000259 130,192,000260 120,777,000242Notes: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 contractingRetrofit capital/CAPCOST-wet FGD-2006.xls.xlsNational 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 Rea g ent 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 B y -product Sale Cost - dr y , $/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 $6,805,000National 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 0.66National 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 Rea g ent 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 B y -product Sale Cost - dr y , $/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 $7,346,000National 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 0.85National 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 $7,235,000National 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 0.64National 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 Rea g ent 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 B y -product Sale Cost - dr y , $/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 $7,775,000National 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