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## T ECHNOLOGY 

## E VALUATION P ROJECT 

### N UMBER 

### 11311

- 000

P REPARED FOR 

## N ATIONAL 

## L IME 

## A SSOCIATION



## JANUARY 2003

P REPARED BY

55 East Monroe Street Chicago, IL 60603

- 5780 USALEGAL NOTICE

This rep ort 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 method s 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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## N ATIONAL 

## L IME 

## A SSOCIATION



### C ONTENTS 

### S ECTION 

### P AGE

Wet FGD

- revision_3

- no merc Project Number 11311

- 000 i 

### 1. FLUE GAS DESULFUR IZATION (FGD) DESCRI PTION

................................ ......................... 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 Proce ss

................................ ................................ ................................ ................................ .......................... 9

1.3 SO 2 Removal Performance

................................ ................................ ................................ ................................ ......... 10

1.3.1 LSFO Process

................................ ................................ ................................ ................................ ....................... 10

1.3.2 MEL Process

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

1.4 Commercial Status

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



### 2. WET FGD PROCESS ADVANTAG ES 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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### S ECTION 

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- 000 ii 3.2.2 SO 2 Removal System

................................ ................................ ................................ ................................ ........... 18

3.2.3 Flue Gas System/Sta ck

................................ ................................ ................................ ................................ ......... 19

3.2.4 By

- Product Treatment and Handling System

................................ ................................ ................................ ....... 20

3.2.5 General Plant Support Systems

................................ ................................ ................................ ............................ 20



### 4. TECHNOLOGY APPLIC ATION CONSTRAINTS

................................ ................................ ... 23

4.1 Unit/Absorber Size

................................ ................................ ................................ ................................ ...................... 23

4.2 Coal Sulfur Content

................................ ................................ ................................ ................................ .................... 23

4.3 Perfo rmance Expectations

................................ ................................ ................................ ................................ ......... 23

4.4 SO 2 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 Co sts

................................ ................................ ................................ ................................ ................ 27

5.2.2 Variable O&M Costs

................................ ................................ ................................ ................................ ............ 28

5.3 Levelized Costs

................................ ................................ ................................ ................................ ............................ 31

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Wet FGD

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- 000 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 Variabl e 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 Cos t/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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- 000 iv R EPORT P REPARED , R EVIEWED , AND A PPROVED BY S ARGENT & L UNDY LLC :

Prepared by: January 17, 200 3

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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- 000 1



## 1. F LUE 

## G AS 

## D ESULFURIZ ATION 

## (FGD) D ES CRIPTION

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 (SO 2 ) reduction from coal

- fired utility boilers. The wet FGD flue g as 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 SO 2 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 (limes tone 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 desig ns, 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 liq uor 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 SO 2 from the gas, facilitating the reaction of the SO 2 with reag ent 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 f

rom 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, SO 2 collection effici ency 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 SO 2 collection efficiency needed and is referred to as liquid

- to

- gas (L/G) ratio. Higher L/G ra tios improve SO 2 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 t his

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- 000 2 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
(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 troubl esome limestone

based "natural oxidation" technology. The oxidized slurry is then recirculated to the spray headers. A portion of the slurry is withdrawn to remove the precipitated gypsum. Typically, this slurry is dewatered in a two

- stage process invol ving 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 corrosiv e 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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- 000 3 1.1 PROCESS CHEMISTRY

1.1.1 LSFO Process

A simplified explanation of the LSFO process (Figure 1) is that the SO 2 absorbed in the recirculated slurry

reacts with dissolved lim estone (CaCO 3 ) in the slurry to form calcium sulfite hemihydrate (CaSO 3 • 1⁄2H 2 O ) according to the following 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 2 is released into the flue gas.

Oxi dation 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 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 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 CaSO 3 • 1⁄2H 2 O crystals produced by older natural oxidation systems. Physical stability of dewate red 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 2 collection efficiency and energy consumption of an LSFO system ca n 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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- 000 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 St ack



### M

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- 000 5 absorbs the acidic SO 2 . This improves absorber performance by i ncreasing solubility of SO 2 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 perc entage 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 l ost 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, glycol ic, 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 ac ids.) 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 so metimes included in designs to allow future addition in case SO 2 collection efficiency is inadequate. Disadvantages of organic acids include additional operating cost, uncertainty in long

- term supply and pricing of DBA by

products, possible contamination of gypsum by

- product, and increased wastewater treatment cost.

1.1.2 MEL Process

In the MEL process, slaked lime, containing calcium hydroxide [Ca(OH) 2 ] and a portion of magnesium hydroxide [Mg(OH) 2 ], is used to react with SO 2 (Figure 2). As with the LSFO pro cess, 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 t he 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, magnesiu m sulfite and magnesium bisulfite [MgSO 3 , Mg(HSO 3 ) 2 ].

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- 000 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

M

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- 000 7 These soluble magnesium salts greatly increase SO 2 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 SO 2 . This improves absorber performance by increasing solubility of SO 2 in the slurry, which allows 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 processes a greater capacity to absorb SO 2 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 SO 2 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) 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

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- 000 8 A portion of recirculated slurry containing absorbed SO 2 in the form of calcium sulfite and magnesium salts is collected in a scoop from recirculat ing 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 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 oxidizer undergoes primary and secondary 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 [95% s maller 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 SO 2 and flu e gas flow rate signals.

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- 000 9 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 agri cultural 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 SO 2 neutralizing value to about 94%

CaO. M agnesium

- 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 SO 2 collec tion 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 hy drogen 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 SO 2 to form magnesium
sulfite, which has the beneficial effects on SO 2 collection effic iency 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 recyc le 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 SO 2 .

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 addit ional cost is assumed for disposing

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- 000 10 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 r ecovery 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 econo mizer outlet, thereby reducing corrosion in the air heater, and lowering sulfuric acid plume potential.



### 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. 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 SO 2 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

W heelabrator, 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 SO 2 basis (high

- sulfur coal case) of this work. This is further verified by recent guarantees offered by FGD vendor s for new unit applications. It should be noted that 98% SO 2 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 onl y be pursued if air permit requirements necessitate very low SO 2 emissions. Note also that these very low SO 2 emissions have been required on some recent air permit applications.

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- 000 11 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% SO 2 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 quali ty 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 con sists 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 tha n 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 2 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% SO 2 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 majo r 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 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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- 000 12

• Chiyoda

• Hamon Research Cottrell

• Mitsubishi

• Wheelabrator Air Pollution Control

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## 2. W ET 

## FGD P ROCESS 

## A DVANTAGES AND 

## D ISAD VANTAGES 

## C OM PARED TO 

## D RY 

## FGD T ECHNOLOGY



### 2.1 P ROCESS 

### A DVANTAGES

The LSFO and MEL technologies have the following advantages when compared wit h 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 98% can be attained.

3. Adequate and commercially viable suppliers offer the technology.

4. Reag ents 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 P ROCESS 

### D ISADVANTAGES

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

1. The LSFO process circulates large quantities of slur ry 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 s ufficiently 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. D ESIGN 

## B ASIS



### 3.1 S PEC IFIC 

### 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 High 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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- 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

* All sulfur in coal is assumed to be converted to SO 2 (no SO 3 ), and that there is no sulfur removed in the either the bott om 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.

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

-

Appalach ian 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 net 2.0 1.3 1.4 1.0 Reagent Limestone Limestone MEL MEL Reagent cost, $/ton 15 1 5 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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- 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 fa ctor, % 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
Discoun t 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 summarize d in Appendix A.



### ** If the LSFO process is to produce a saleable grade gypsum, 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 descriptions present important features of the components and assumptions used in this study:

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- 000 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 conveyor system to transfer limestone to a 16

- hour capacity day bin. The limestone day bin and a gravimetric feeder supply the limesto ne 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

classificat ion 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 stor age 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 sulfu ric acid feed pumps is also provided.

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- 000 18 3.2.2 SO 2 Removal System

3.2.2.1 LSFO Process

One 100% capacity absorber is provided to achieve 98% SO 2 removal efficiency. The absorber is a vertical

open

- spray tower, with countercurrent contact between the flue gas and scrubber slurry. To achieve 98% SO 2 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% oxidatio n 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 oxidized air compressors. The bleed slurry from the absorber is pumped to the dewatering area. Makeup limestone slurry is

added to the react ion 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% SO 2 removal efficiency. The absorber is a vertical

open

- spray tower, w ith countercurrent contact between the flue gas and scrubber slurry. To achieve 98% SO 2 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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- 000 19 L/G designs. The absorber is constructed of 317 LMN steel or equivalent. The chloride lev els 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 CaSO 3 to CaSO 4 . 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

CaSO 3 to C aSO 4 , a 3:1 stoichiometric ratio of oxygen to removed SO 2 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 dewate ring 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 overal l 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 Sy stem/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 accomm odated by the booster ID fans. The booster ID fan is sized to provide an additional 10" H 2 O (9"

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- 000 20 operating) and 9” H 2 O (8" operating) pressure drop through the absorber for high

- and low

- sulfur applications, respectively. The inlet ductwork to the absorbe r 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” H 2 O (7” operating).

3.2.4 By

- Product Treatment and Handling System

The by

- product system will be simi lar 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 th e 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 filter s 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 conveyor s 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 additio ns to the plant seal water system, instrument air system, makeup water system, distributed control system (DCS), and electrical auxiliary power

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- 000 21 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 E QUIPMENT U SED IN E ACH S UBS YSTEM 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

Sl urry 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)

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:

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- 000 22 1 x 100% module with reaction tank and slurry trough 3 recirculation pumps (2 opera ting, 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

Makeu p 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 sewer s Sanitary systems Fire protection

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## 4. T ECHNOLOGY 

## A PPLIC ATION 

## 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 n umber 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 demon strated and available from industry suppliers.



### 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 are

installed on medium

- to high

- sulfur fuels (>2%); however, a number of systems are al so installed on low

sulfur fuels (<1%) in the western areas of the United States.



### 4.3 P ERFORMANCE 

### E XPECTATIONS

The SO 2 removal guarantees of up to 98% are available from the system suppliers and have been demonstrated in commercial applications.



### 4.4 SO 2 

### R EDUCTION

Suppliers of FGD systems have guaranteed SO 2 reduction efficiency up to 98% with the inlet SO 2 concentration up to 10.0 lb/MBtu. To the best of our knowledge, the lowest outlet SO 2 loading guaranteed by a supplier of FGD systems is 0.06 lb/MBtu.

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### 4.5 R EAGENT 

### U TILIZATION

Reagent utilization is the percentage of reagent fed to the FGD system that reacts with SO 2 . Reagent utilization in a wet FGD system is typically very close to the required theoretical level when considered on a removed

SO 2 basis. Limestone re agent 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 SO 2 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 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 manu facturing, or for agricultural use. The by

- product quality requirements are much more stringent for wallboard manufacturing compared to the other uses.



### 4.7 E NERGY 

### C ONSUMPTION

The process energy consumption results primarily from the booster ID fan power requi red 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 (high

sulfur case) than for th e LSFO process and the power requirement for the booster ID fan is 22% lower (high

sulfur case) than for the LSFO process. Overall, the MEL process will require approximately 0.6% less auxiliary (% of MW net ) power for high

- sulfur coal and approximately 0. 3% less auxiliary power auxiliary (%

of MW net ) for low

- sulfur coal compared to the LSFO process.

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### 4.8 R ETROFIT 

### V ERSUS 

### N EW 

### U NITS

In new unit and retrofit applications, wet FGD technology is typically installed between the electrostatic precipitator/baghouse ou tlet 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 o vercome 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 veloc ity 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. C OSTS 

## A NA LYSES



### 5.1 C APITAL 

### C OSTS

Estimated capi tal 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 th e 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 s pare 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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- 000 27 Salient features of each capital cost estimate are described in Section 3 of this report.

Based on our experience and the leve l 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 s pecialty 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 continge ncy, 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 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 o f 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 th e 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, th e required maintenance material, and the labor required to perform the maintenance.

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Variable O&M costs developed for each FGD system include the cost of all consumables, including reagent (limestone or lime), water, and power requiremen ts. 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 p ower 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 ty pes 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 dispo sal.

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 val ues,

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 econom ic 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 (inclu ding 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 wou ld be the market

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- 000 29 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 tho se 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:

T ABLE 5.2.2

- 1 R ECENT T OTAL S POT M ARKET E LECTRICITY C OSTS $/MWh
Region On Pea k 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:

E CAR 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 p eak 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 electrici ty 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 f easibility 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 a pplications, these costs tend toward the higher end of the range.

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– 3 and Exhibit 5

– 4 represent the fixed and variable O&M costs for new and retrofit application s 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 techn ology, respectively.



### 5.3 L EVELIZED 

### C OSTS

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 in surance. 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.

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- 1 APPENDIX A

A UXILIARY P OWER S UMMARY

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:

T ABLE A

- 1 E STIMATED FGD S YSTEM P RESSURE D ROP LSFO High

Sulfur LSFO Low

Sulfur MEL High

Sulfur MEL Low

Sulfur

Ductwork to absorber (“w.c.)* 1 1 1 1

Absorber inlet expansion (“w.c.) 0.5 0.5 0.5 0.5

Ab sorber 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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- 2 T ABLE A

- 2 P OWER R EQUIREMENTS WITH F LUE G AS H ANDLING AND O THER FGD S UBSYSTEMS LSFO High

Sulfur LSFO Low

Sulfur MEL High

Sulfur MEL Low

Sulfur Reagent Preparation (kW) 530 225 330 140

SO 2 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

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- 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 slurr y 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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- 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

h ydroxide [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



### S O + H O HS O + H HS O + H + MgS O Mg(HS O ) 2 2 3

3 3 3 2 - +

- +



### Mg(HS O) + Ca(OH ) CaS O 1⁄2 H O (solid) + MgS O (aqueous ) + 1-1/2 H O 2 Mg(HS O) + Mg(OH ) 2 MgS O + + 2 H O

3 2 2

3 2

3 2

3 2 2

3 2 

### Mg(HS O )

3 2 

### 102 S O + 94 Ca(OH ) + 6 Mg(OH) 94 CaS O 1⁄2 H O (solid) + 4 MgS O (aq.) + 2

2 2 2

3 2

3 

### 51 H O

Mg(HS O ) (aq.) 2

3 2



### STEP 1: ABSORPTION

STEP 2: PRECIPI TA

TION AND REGENER ATION: OVERALL:



### SO 2 

### S O ABSORBED BY MgSO 2

3 

### Mg(HS O )

3 2



### MgS O in solution

3 

### Flue gas with SO 2 

### Cleaned Gas

To Forced

Oxidation & Gypsum Production

Slaked Lime 94 moles Ca(OH) 6 moles Mg(OH)

2

2 

### CaS O 1⁄2 H O

3 2 

### M

The overall reaction in the absorber of lime with S O 2 , for a magnesium

- enhanced lime with 6 moles of magnesium hydroxide and 94 moles of calcium hydroxide, is shown in reaction (1):

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- 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

Alth ough reaction (1) shows the overall reaction, SO 2 does not react directly with calcium or magnesium hydroxide. SO 2 capture with lime occurs in two steps. In the first step, SO 2 is absorbed into the liquid phase of slurry recirculated from the recycle tan k through 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 slurr y 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 sulfurous 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 o f 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 rea cts with a smaller portion of

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- 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 c alled 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 just contacted the flue gas is collected in a sc oop, 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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- 4 the absorber via the scoop contains ca lcium 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 oxidation tank where it is contacted with comp ressed 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 a fter 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 n ecessary 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 ove rall 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 utilizati on 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 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.xlsNational 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.xlsNational 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 W ater Requirement, gpm 660 660 12 W ater 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 0.44National 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 W ater Requirement, gpm 660 660 12 W ater 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 0.56National 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 0.43National 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 W ater Requirement, gpm 660 660 12 W ater 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