# Chapter 2 Selective Catalytic Reduction

John L. Sorrels Air Economics Group Health and Environmental Impacts Division Office of Air Quality Planning and Standards U.S. Environmental Protection Agency Research Triangle Park, NC 27711

David D. Randall , Karen S. Schaffner, Carrie Richardson Fry RTI International Research Triangle Park , NC 277 09

May 201 6Chapter 2

– Selective Catalytic Reduction

CONTENTS

2. Selective Catalytic Reduction

................................ ................................ ................................ ..... 2 2.1 Introduction

................................ ................................ ................................ ...................... 2 2.2 Process Description

................................ ................................ ................................ ........ 10 2.2.1 Reduction Chemistry, Reagents, and Catalyst

................................ ................... 11 2.2.2 SCR Performance Parameters

................................ ................................ ............ 16 2.2.3 SCR System Configurations

................................ ................................ .............. 29 2.2.4 SCR System Primary Equipment

................................ ................................ ....... 35 2.2.5 SCR System Auxiliary Equipment

................................ ................................ .... 42 2.2.6 Other Consideratio ns

................................ ................................ ......................... 45 2.3 Design Parameters

................................ ................................ ................................ ......... 50 2.3.1 Boiler Heat Input

................................ ................................ ................................ 51 2.3.2 Heat Rate Factor

................................ ................................ ................................ 52 2.3.3 System Capacity Factor

................................ ................................ ..................... 52 2.3.4 Inlet NO x and Stack NO x ................................ ................................ ................... 54 2.3.5 NO x Removal Efficiency

................................ ................................ .................... 54 2.3.6 NO x Removal Rates

................................ ................................ ........................... 54 2.3.7 Actual and Normalized Stoichiometric Ratios

................................ .................. 55 2.3.8 Flue Gas Flow Rate

................................ ................................ ............................ 55 2.3.9 Space Velocity and Area Velocity

................................ ................................ ..... 56 2.3.10 Theoretical NO x Removal Efficiency

................................ ................................ 57 2.3.11 Catalyst Volume

................................ ................................ ................................ . 58 2.3.12 SCR Reactor Dimensions

................................ ................................ .................. 59 2.3.13 Estimating Reagent Consumption and Tank Size

................................ .............. 61 2.4 Cost Analysis

................................ ................................ ................................ ................. 62 2.4.1 Total Capital Investment

................................ ................................ .................... 64 2.4.2 Total Annual Costs

................................ ................................ ............................ 72 2.5 Example Problem #1

– Utility Boiler

................................ ................................ ............ 79 2.5.1 Design Parameter Example #1

................................ ................................ ........... 80 2.5.2 Cost Estimation Example

................................ ................................ ................... 84 2.6 Example Problem #2

– Industrial Boiler

................................ ................................ ....... 87 2.6.1 Design Parameter Example #2

................................ ................................ ........... 88 2.6.2 Cost Estimation Example #2

................................ ................................ .............. 91 References

................................ ................................ ................................ ............................... 96Chapter 2

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- 2 

## 2. SELECTIVE CATALYT IC REDUCTION

2.1 Introduction

Selective catalytic r eduction (SCR) has be en applied to stationary source fossi l fuel

– f ired combustion units for emission control since the early 1970s and is currently being used in Japan, Europe, the United States , and other countries

. In the U.S. alone, more than 1,000 SCR systems have been installed on a wide variety of sources i n many different industries, including utility and industrial boilers, process heaters, gas turbines, internal combustion engines, chemical plants

, and steel mills [ 1 ]. Other sources include fluid catalytic cracking units (FCCUs) , ethylene cracker furnaces , nitric acid plants, catalyst manufacturing processes, nitrogen fixation processes, and solid/liquid or gas waste incinerators [ 2, 3 ]

. In the U.S., SCR has been installed on m o re than 300 coal

- fired power plants ranging in size from <100 MWe to 1,400 MWe [ 1, 4 ]. Other combustion sources with large numbers of SCR retrofits include more than 50 gas

- fired utility boilers ranging in size from 147 MWe to 750 MWe, more than 50 industrial boilers and process heaters (both field

- erected and packaged units), and mo re than 650 combined cycle gas turbines [ 1 ]. SCR can be applied as a stand

- alone NO x control or with other technologies , including selective non

- catalytic reduction (SNCR) 1 and combustion controls such as low NO x burner ( LNB ) and flue gas recirculation ( FG R ) [ 2 ]

. SCR is typically implemented on stationary source combustion units requiring a higher level of NO x reduction than achievable by selective noncatalytic reduction ( SNCR ) or combustion controls. Theoretically, SCR systems can be designed for NO x removal efficiencies
up close to 100 percent (%). In practice, c ommercial coal

- , oil

- , and natural gas

– fired SCR systems are often designed to meet control targets of over 90%. However, the reduction may be less than 90% when SCR follow s other NOx control s such as LNB or FGR that achieve relatively low emissions on their own. The outlet concentration from SCR on a utility boiler rarely is less than 0.04 lb/MMBtu [1]. 2 In comparison, SNCR units typicall y achieve approximately 25 to 75 % reduction efficiencie s [ 5 ]

. Either ammonia or urea may be used as the NO x reduction reagent in SCR systems. Urea is generally converted to ammonia before injection. Results of a survey of electric utilities that operate SCR systems indicated that about 80 percent use ammonia ( anhydrous and aqueous) , and the remainder use urea [ 4 ]. A survey of coal

- fired power plants that control NO x emissions using either SCR or SNCR found anhydrous ammonia use exceeds aqueous ammonia use by a ratio of 3 to 1. Nearly half of these survey respon dents also indicated that price is their primary consideration in the choice of reagent; safety is the primary consideration for about 25 percent of the operators [ 6 ].
SCR capital costs vary by the type of unit controlled

, the fuel type , the inlet NOx level, the outlet NO x design level, and reactor arrangement

. Capital costs also rose between 2000 and 2010 (at least for utility boiler applications) , even after scaling all data to 20 11 dollars. For a

1 A hybrid SNCR/SCR system was demonstrated at the AES Greenidge Power Plant in 2006. However, no hybrid SNCR/SCR systems are currently known to be operating as of February 2016.
2 Data in the Clean Air Markets Division (CAMD) database also suggest SCR units rarely achieve emissions less than 0.04 lb/MMBtu.Chapter 2

– Selective Catalytic Reduction

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- 3 small number of early SC R retrofits on utility boilers prior to 2000

, the average costs were about $100 /k W, in 20 11 dollars , and there was little scatter in the data . From 2000 to 200 7 , the SCR costs for 32 utility boilers ranged from about $ 10 0/k W to $2 75 /k W (2011$) , and a slight economy of scale was evident (i.e., using a regression equation, costs ranged from about $200/kW for a 200 MW unit to $160/kW for an 800 MW unit) . For 2008 to 201 1 , the average SCR costs exhibited great variability and again a modest economy of scale was evident (i.e., about $300/kW for a 200 MW unit to $250/kW for an 800 MW unit ; 2011$ ) . For eight utility boilers either installed in 2012 or projected to be installed by 2014, the SCR costs ranged from about $270/kW to $570/kW, in 2011$; generating capacity for these units ranged from 400 MW to 800 MW [ 7 b ]. Typical operation and maintenance costs are approximately 0.1 cents per kilowatt

- hour (kWh) [7 a , 8 ]

. Table 2. 1 a provides capital cost estimates for electric utility boilers, and Table 2.1b presents capital cost estimates for SCR applications of various sizes in several other industry source categories.

T he procedures for estimating costs presented in this report are based on cost data for SCR retrofit s on existing coal

- , oil

- , and gas

- fired boilers for electric generating units larger than 25 MWe (approximately 250 MMBt u/hr) . Thus, this report ’ s procedure estimat es costs for typical retrofits of such boilers. The methodology for utility boilers also ha s been extended to large industrial boilers by modifying the capital cost equations and power consumption (electricity cost) equations to use the heat input capacity of the boiler instead of electric generating capacity. 3 The procedures to estimate capita l costs are not directly applicable to sources other than utility and industrial boilers. Procedures to estimate annual costing elements other than power consumption are the same for SCR units in any application. The cost for SCR as part of a new plant oft en is likely to be less than the cost for retrofit plant

. Appropriate factors to estimate the cost of a new plant SCR have been included. In addition, the cost procedures in this report reflect individual SCR applications. Retrofitting multiple boilers w ith SCR can allow for some economies of scale for installation, thus yielding some reduction in capital costs per SCR application. The cost methodology incorporates certain approximations; consequently, it should be us ed to develop study

- level accuracy (± 3 0%) cost estimates of SCR applications. Such accuracy in the cost methodology is consistent with the accuracy of the cost estimates for the other control measures found in this Cost Manual as stated in Section 1

. In the cement industry, pilot tests in the 1970s and 1990s showed SCR could be a feasible control technology for cement kilns. Building on that experience, SCRs were first installed in Europe in 2001. Today, SCR ha s been successfully implemented at seven European cement plants in Solnhofer, German y (operated from 2001 until 2006), Bergamo, Italy (2006), Sarchi, Italy (2007), Mergelstetten, Germany (2010), Rohrdorf, Germany (2011), Mannersdorf, Austria (2012), and Rezatto, Italy (2015). [94, 98, 99]. As of 2015, t here is only one cement plant in the U.S. has installed an SCR. This SCR began operation in 2013 and is installed after an electrostatic precipitator. The control efficiency for the system is reported to be about 80%, which is consistent with SCR applications on European kilns. SCRs have not seen widespread use in the U.S. cement industry mainly due to industry concerns regarding potential problems caused by high

- dust levels and catalyst deactivation by high SO 3 concentrations from pyritic sulfur found in the raw materials used by U.S. cement plants. The SO 3 could react with calcium

3 “Industrial” boilers as a term used in the Control Cost Manual is meant to cover not only industrial but also commercial and institutional (or ICI) boilers, unless noted otherwise.Chapter 2

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- 4 oxide in the flue gas to form calcium sulfate and with ammonia to form ammonium bisulfate. The calcium sulfate could deactivate the catalyst, while the ammonium bisulfate could cause catalyst plugging. The re have b een concerns expressed about the potential for catalyst poisoning by sodium, potassium, and arsenic trioxide. Finally, other concern s expressed are that dioxins and furans may form in the SCR due to combustion gases remaining at temperatures between 450°F and 750°F. These and other concerns regarding the implementation of SCR to the cement industry are discussed in detail in “Alternative Control Techniques Document Update

– NOx Emissions from New Cement Kilns” [94]. Due to the small number of SCRs installed at cement plants, information on capital and operating costs for SCRs at cement plants is limited. The installation and operating costs for the SCR installed at the U.S. plant in 2013 are not publicly available at this time. In general, we expect the capi tal and operating costs would be higher than for low

- dust applications due to the need to install catalyst cleaning equipment for SCR systems installed in high

- dust configurations and for heating the flue gas in low

- dust, tail

- end configurations.Chapter 2

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- 5 Table 2.1a: Summary of SCR Cost Data for Utility Boilers

Source Category Unit Size Fuel Type Capital Cost $ Year Comments Reference Min Avg Max

Electric Generating Units NA a NA $55/kW $140/kW <2000$ b Retrofit costs. [9]

~300

- 1,400 MW NA ~$70/kW ~$120/kW <2000$ b Retrofit costs. Six boilers. No economy of scale. [9]

150

– 1,000 MW Coal $80/kWnet $160/kWnet 2002$ Retrofit costs. Author of referenced document scaled original costs to 2002 dollars. More than 20 boilers. Little to no economy of scale. [10]

NA Coal $60/kW $100kW $200/kW <2004$ b Retrofit costs [11]
<300 MW Coal $167/kW $186/kW <2004$ Costs for 26 boilers. [12]

301

– 600 MW Coal $148/kW $192/kW <2004$ Costs for 15 boilers. [12]

601

– 900 MW Coal $124/kW $221/kW <2004$ Costs for 22 boilers. [12]
>900 MW Coal $118/kW $195/kW <2004$ Costs for 9 boilers. [12]

100

– 399 MW Coal $70/kW $123/kW ~$175/kW <2004$ b Costs for 5 boilers. [13]

400

– 599 MW Coal $73/kW $103/kW ~$160/kW <2004$ b Costs for 8 boilers. [13]

600

– 899 MW Coal $56/kW $81/kW ~$100/kW <2004$ b Costs for 9 boilers. [13]
>900 MW Coal ~$80/kW $117/kW ~$190/kW <2004$ b Costs for 10 boilers. [13]
191 MW Coal $149/kW 2006$ Retrofit costs. [14]
~100 MW

- ~800MW NA ~$125/kW $275/kW ~$440/kW 2008$ Retrofit costs for 15 boilers installed in 2008 to 2010 . Most costs between $200/kW and $350/kW. Slight economy of scale

— regression average about $340/kW for 100 MW to $250/kW for 800 MW. [7 a ]

~400 MW to ~800 MW NA ~$270/kW ~$420/kW ~560/kW 2011$ Retrofit costs for 8 boilers either installed in 2012 or projected to be installed by 2014. [7b]

a Not Available.

b Year of reference.Chapter 2

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- 6 Table 2.1 b : Summary of SCR Cost Data for Miscellaneous Industrial Sources

Source Category Unit

Size Fuel

Type Capital Cost: average

(range) $ Year Actual , Vendor Quote, or Estimated? Comments Reference

Industrial

Commerci al Boilers 350

MMBtu Coal NA ($10,000

– $15,000/MMBtu/hr) 1999$ Estimated Retrofit costs. Authors of referenced document estimated the low end of the range assuming a cost of about $100/kW for a 100 MW (1000 MMBtu/hr) utility boiler and assuming that economies of scale would be greater for utility boilers than for industrial boilers (so that the cost for a 350 MMBtu/hr industrial boiler would be comparable to or greater than the cost for a 1000 MMBtu/hr utility boiler on a $/MMBtu basis). [15]

100

– 1,000 MMBtu/ hr Coal NA ($7,300

– $14,600/MMBtu/hr) 1999$ Estimated Retrofit costs. Generally costs available for one boiler with each type of fuel. Authors of referenced document estimated costs for other sizes assuming ratio of small

- to

- large $/MMBtu costs are related to ratio of large to small heat inputs raised to the 0.3 power. [1 6 ]

100

– 1,000 MMBtu/ hr Oil N A ($5,550

– $11,100/MMBtu/hr) 1999$ Estimated [1 6 ]

100

– 1,000 MMBtu/ hr Gas NA ($4,010

– $8,010/MMBtu/hr) 1999$ Estimated [1 6 ]

100 MMBtu/ hr Gas NA ($7,500/MMBtu/hr) 1999$ b V e ndor Cited source in reference [15] is an unpublished letter from a vendor. [1 5 ]

350 MMBtu Oil, Gas,

or Wood NA ($4,000

–

$6,000/MMBtu/hr) 1999$ Estimated [ 17 ]

57 MMBtu/ hr Wood NA (>$560,000 and $9,500/MMBtu/hr) 1999$ c Actual / Estimate Costs for a new boiler. [1 5 ]

321 MMBtu/ hr Wood NA ($1,980/MMBtu/hr) 2006$ Likely Estimated [1 8 ]Chapter 2

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- 7 Source Category Unit

Size Fuel

Type Capital Cost: average

(range) $ Year Actual , Vendor Quote, or Estimated? Comments Reference

Petroleum Refining

– Steam Boilers 650

MM B tu/ hr Gas or

refinery fuel gas NA ($ 3,100

–

$ 25,800/MMBtu ) 2004$ c Estimated Retrofit costs. Equipment costs based on range of costs found in literature search (references were not provided). Installation costs estimated using factors from the Control Cost Manual for thermal and catalytic incinerators. [1 9 ]

Petroleum Refining

– Process Heaters 350

MMBtu/ hr Gas/refin

ery fuel gas NA ($ 3,100

–

$ 25,800/MMBtu ) 2004$ c Estimated Same comment as above. [1 9 ]

350 MMBtu/ hr Refinery

oil NA ($ 3,100

–

$ 25,800/MMBtu ) 2004$ c Estimated Same comment as above. [1 9 ]

10 MMBtu/ hr Gas or

refinery fuel gas/NG combo $19 ,200/MMBtu ($ 12,000

–

$ 26,500/MMBtu ) 1999 b Vendor / Estimated Costs are based primarily on quotes from two

vendors (and additional discussions). Authors of the referenced report added costs for fan, motor, and ductwork costs based on procedures in the Control Cost Manual. [ 20 ]

50 MMBtu/ hr Gas or

refinery fuel gas/NG combo $ 5,140/MMBtu ($ 4,020

–

$ 6,280/MMBtu ) 1999 b Vendor / Estimated Same comment as above. [ 20 ]

75 MMBtu/ hr Gas or

refinery fuel gas/NG combo $ 4,190/MMBtu ($ 3,440

–

$ 4,950/MMBtu ) 1999 b Vendor / Estimated Same comment as above. [ 20 ]

150 MMBtu/ hr Gas or

refinery fuel gas/NG combo $ 2,730/MMBtu ($ 2,570

–

$ 2,880/MMBtu ) 1999 b Vendor / Estimated Same comment as above. [ 20 ]

350 MMBtu/ hr Gas or

refinery fuel gas/NG combo $ 1,550/MMBtu ($ 1,520

–

$ 1,570/MMBtu ) 1999 b Vendor / Estimated Same comment as above. [ 20 ]Chapter 2

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- 8 Source Category Unit

Size Fuel

Type Capital Cost: average

(range) $ Year Actual , Vendor Quote, or Estimated? Comments Reference

68 MMBtu/ hr (Two 32 MMBtu/ hr) Refinery

fuel gas NA ($ 22,100/MMBtu ) 1991 Actual Retrofit costs. [1 5 ]

Petroleum Refining

– FCCU 70,000

barrels/s t ream day (bbl/stre am day) NA NA ($9. 0 million) 2004$ c Vendor Estimated cost by vendor (for 90% reduction)

. [ 3 ]

27,000 bbl/strea m day NA NA ($8

- $12 million) 2009 Estimated [21]

<20,000

- >100,00 0 bbl/strea m day NA NA (order of magnitude range; low end higher than two entries above) 2005 to 2010 Actual Costs reported by 6 petroleum refining companies for 7 FCCUs in responses to EPA ICR. One new, 6 retrofits. [22]

NA NA NA ($20 million) 2006 Actual Approximate average cost for SCR retrofits at several refineries [23]

Portland Cement (dry kilns) 1.0 9

million short tpy clinker NA NA ($ 6.9 per short ton clinker) 2006 a Estimated Retrofit cost. Estimate based primarily on SCR procedures for boilers in fifth edition of the Control Cost Manual. Clinker capacity obtained from the second reference. [ 24,25 ]

1.13 million short tpy clinker NA NA ($ 5.9 per short ton clinker) 2006 a Estimated Same comment as above. [ 24,25 ]

2.16 million short tpy clinker NA NA ($3.9 per short ton clinker) 2006 a Estimated Same comment as above. [24,25]Chapter 2

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- 9 Source Category Unit

Size Fuel

Type Capital Cost: average

(range) $ Year Actual , Vendor Quote, or Estimated? Comments Reference

1.4 million short tpy clinker NA NA ($5.9 per short ton clinker) 2004 Not clear Retrofit cost for European kiln. Cost in euros converted to dollars assuming a ratio of $1.3/euro. [26]

1.055 million tpy clinker NA NA ($4. 4 per short ton clinker ) 2004 Estimated Cost for new kiln. [ 27 ]

1.095 million short tpy clinker NA NA ($ 4.4 per short ton clinker ) 20 11 Estimated Cost for new kiln. Cost based on quote for the SCR equipment, and standard installation factors from the Control Cost Manual for other types of control devices. [ 28 ]

Portland Cement (wet kilns) 0.3

million short tpy clinker NA NA ($ 17.5 per short ton clinker) 2006 a Estimated Retrofit costs for 4 kilns. Rated clinker production capacity obtained from the second reference. [ 24,29 ]

0.320 million short tpy clinker NA ($15.6

- $16.6 per short ton clinker) 2006 a Estimated Retrofit costs for 3 kilns. Rated clinker production capacity obtained from second reference. [24,25]

Gas Turbine, Simple Cycle NA Gas NA ($50

- $70/kW) 1999$ a Vendor Retrofit costs. [1 5 ]

80 MW Gas NA ($51/kW) 1999$ a Vendor Retrofit cost, excluding balance of plant costs. [1 5 ]

2 MW Gas NA ($ 237/kW ) 1999$ a Vendor Retrofit cost. [1 5 ]

12 MW Gas NA ($167/kW) 1999$ a Vendor Retrofit cost. [1 5 ]

Internal Combusti on Engine 1,800 hp Diesel (No. 2 fuel oil) NA ($0.18 million) 1994 Actual New cost [1 5 ]

a Year of reference.

b Year analysis was conducted (assumed vendor contacts were made that year).

c Commission year of the SCR.Chapter 2

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- 10 

### 2.2 Process Description

Like SNCR, the SCR process is based on the chemical reduction of the NO x molecule. The primary difference between SNCR and SCR is that SCR employs a metal

- based catalyst with activated sites to increase the rate of the reduction reaction. The primary components of the SCR include the ammonia storage and delivery system, ammoni a injection grid, and the catalyst
reactor [ 2 ]

. A nitrogen

- based reducing agent (reagent), such as ammonia or urea

- derived ammonia , is injected into the post

- combustion flue gas. The reagent reacts selectively with the flue gas NO x within a specific temperature range and in the presence of the catalyst and oxygen to reduce the NO x into molecular nitrogen ( N 2 ) and water vapor ( H 2 O ).
The use of a catalyst results in two primary advantages of the SCR process over SNCR. The main advanta ge is the higher NO x reduction efficiency. In addition, SCR reactions occur within a lower and broader temperature range. However, the decrease in reaction temperature and increase in efficiency is accompanied by a significant increase in capital and opera ting costs. The capital cost increase is mainly due to the large volumes of catalyst required for the reduction reaction

. Operating costs for SCR consist mostly of replacement catalyst and ammonia reagent costs, and while historically, the catalyst replace ment cost has been the largest cost, the reagent
cost has becom e the most substantial portion of operating costs for most SCR [ 7 b ]

. 4
Figure 2.1 shows a simplified process flow schematic for SCR. Reagent is injected into the flue gas downstream of the combu stion unit and economizer through an injection grid mounted in the ductwork. The reagent is generally diluted with compressed air or steam to aid in injection. The reagent mixes with the flue gas

, and both components enter a reactor chamber containing the catalyst. As the hot flue gas and reagent diffuse through the catalyst and contact activated catalyst sites, NO x in the flue gas chemically reduces to nitrogen and water. The heat of the flue gas provides energy for the reaction. The nitrogen, water vapor, and any other flue gas constituents then flow out of the SCR reactor. More detail on the SCR process and equipment is provided in the following sections.

There are several different locations downstream of the co mbustion unit where SCR
sys tems can be inst alled. Flue gas temperature and constituents vary with the location of the SCR reactor chamber. SCR reactors located upstream of the particulate control device and the air heater (“high

- dust ” configuration) have higher temperatures and higher levels of par ticulate matter. An SCR reactor located downstream of the air heater, particulate control devices

, and flue gas desulfurization (FGD) system ( “low

- dust” or “tail

- end” configuration) i s essentially dust

- and sulfur

- free but its temperature is generally below the acceptable range. In this case, reheating of the flue gas may be required

, which significantly increases the SCR operational costs. Section 2.2.3 discusses the various SCR system configurations.

4 Several cost analyses in recent years have shown the largest operating cost is for reagent usage rather than for catalyst costs. For example, for the Navajo Generating Station in Arizona, a 2010 BART analysis report on an 812 MW gross coal

- fired unit estimates annual operating costs for ammonia reagent of $1,035,000 (based on $465/ton) and for catalyst replacement of $672,000 (based on $8,000/m 3 ) [30].Chapter 2

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- 11 Figure 2.1: SCR Process Flow Diagram [ 31, 32 ]

2. 2.1 Reduction Chemistry, Reagents , and Catalyst
The reducing agent employed by the majority of SCR systems is gas

- phase ammonia ( NH 3 ) because it readily penetrates the catalyst pores. The ammonia, either in anhydrous or aqueous form, is vaporized before i njection by a vaporizer. Within the appropriate temperature range, t he gas

- phase ammonia then decom poses into free radicals

, including NH 3 and NH 2 . After a series of reactions, the ammonia radicals come into contact with the NO x and reduce it to N 2 and H 2 O

. Since NO x includes both NO and NO 2 , the overall reactions with ammonia are as follows:

(2.1 a )

(2.1b)

The equations indicate that one mole of NH 3 is required to remove one mole of NO and two moles of N H 3 are required to remove one mole of NO 2 . However, Equation 2.1a is theChapter 2

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2

- 12 predominant reaction because 90 to 95% of NO x in flue gas from combustion units is NO. Hence, about one mole of NH 3 is required to remove one mole of NO x . The catalyst lowers the required activation energy for the reduction reaction and increases the reaction rate. In the catalytic reaction, activated sites on the catalyst rapidly adsorb ammo nia and gas

- phase NO to form an activated complex. The catalytic reaction, repre sented by Equation s 2.1 a and 2.1b

, result s in nitrogen and water, which are then desorbed to the flue gas. The site at which the reaction occurs is then reactivated via oxidation.
The high temperature of the flue gas converts the ammonia t o free radicals and provides the activation energy for the reaction. The reaction also requires excess oxygen, typically 2

– 4 %

, to achieve completion. NO x reduction with ammonia is exothermic, resulting in the release of heat. However, because the NO x conce ntration in the flue gas at the inlet of the SCR is typically 0.01

– 0.02 % by volume, the amount of heat released is correspondingly small. Thermodynamic equilibrium is not a limiting factor in NO x reduc tion if the flue gas is within the required temperature range [ 33 ]

. Reagent
The SCR system can use either aqueous or anhydrous ammonia for the reduction reaction , and some plants use urea

- to

- ammonia reagent systems where aqueous ammonia is produced onsite (often called onsite urea

- derived ammonia production or “ammonia

- on

demand” ) . Anhydrous ammonia is nearly 100% pure ammo nia. It is a gas at normal atmo spheric temperature ; therefore, it must be transported and stored under pressure. Anhydrous ammonia is classified as a hazardous material and often requires special permits as well as additional procedures for transportation , handling and storage.
SCR applications using aqueous ammonia generally transport and store it at a

con centration of 29.4% ammonia in water, alt hough some appl ications use a 19% solu tion [ 33 ]

. The use of aqueous ammonia reduces transport and storage problems related to safety . In addition, certain locations may not require permits for concentrations less than 28%. Aqueous ammonia

, however, requires more storage capacity than anhydrous ammon ia and also requires shipping costs for the water solvent in the solution

. Although the 2 9.4% solu tion has substantial vapor pressure at normal air temperatu res, a vaporizer is generally required to provide sufficient

ammo nia vapor to the SCR system. Table 2.2 gives the properties of anhydrous ammonia and the properties of a 29.4% aqueous ammonia solution.
The type of reagent used affects both the capital costs and annual costs. Anhydrous ammonia typically has the lowest capital and operating costs, excluding highly site

- dependent permitting and risk management planning and implementation costs. Urea systems have the highest capital costs due to the complexity of the processing eq uipment. Aqueous ammonia systems tend to have the highest operating costs, primarily because of the cost for transportation. Urea systems have the highest energy consumption costs because the energy needed to hydrolyze or decompose urea tends to be higher than the energy needed to vaporize aqueous ammonia. Although the price per ton of anhydrous ammonia is higher than the price per ton of urea, the cost per ton of NO x removed is higher for urea due to urea’s much higher molecular weight. For example, one SC R supplier estimated capital costs for a 130 lb/hr ammonia system to be $280,000 for anhydrous ammonia, $402,0 0 0 for 19% aqueous ammonia, and $750,000 for urea [34 ]

. Another reference reported that the equipment cost for urea is generally twice theChapter 2

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2

- 13 equipme n t cost for anhydrous ammonia [35 ]. According to one reference, the total SCR system cost is 2

- 5% higher when using a urea reagent system instead of an anhydrous ammonia system [10 ]. Relative to anhydrous ammonia, o ne reference estimated annual operating c osts for 1 9% aqueous ammonia are 50% high er , costs for 29% aqueous ammonia are 33 percent higher, and costs for urea are 25% higher [36 ]

. Another reference stated that as a general rule, operating costs for urea systems are about 50 percent more than the o perating costs for anhydrous ammonia [35 ]. One reference estimated energy costs for an unspecified application to be $167,000 for a urea syste m, $73,000 to $117,000 for aque ous ammonia systems, and $ 16,000 for anhydrous ammonia [37 ]

. This presentation is valid for anhydrous or aqueous ammonia ; the capital cost procedures are based on the typical mix of systems actually in operation, while the procedures for estimating annual costs apply to any ammonia system (the example s in section 2.5 illustrate the proc edures for a system using 29% aqueous am monia as the reagent )

. Table 2.2 : Ammon ia Reagent Properties

Property Anhydrous Ammonia [ 38,39 ] Aqueous Ammonia Liquid or gas at normal air temperature Liquid at high pressure; gas at atmospheric pressure Liquid

Concentration of reagent normally supplied 99.5% (by weight) 29.4% (by weight of NH 3 )

Molecular weight of reagent 17.03 17.03 (as NH 3 )
Ratio of ammonia to solution 99.5% (by weight of NH 3 ) 29.4% (by weight of NH 3 ) Density of liquid at 60°F 5.1 lb/gal 7. 5 lb/gal Vapor pressure at 80°F 153 psia 14.6 psia [ 39 , p. 3] Flammability limits in air 16

– 25% NH 3 (by volume) 16

– 25% NH 3 (by volume)
Short

- term exposure limit 35 ppm 35 ppm
Odor Pungent odor at 5 ppm or more Pungent odor at 5 ppm or more Acceptable materials for storage Steel tank, rated for at least 250 psig pressure (no copper or copper

- based alloys, etc.) Steel tank, rated for at least 25 psig pressure (no copper or copper

based alloys, etc.) Catalyst
SCR catalysts are composed of active metals or ceramics with a highly porous structure. Within the pores of the catalyst are activated sites. These sites have an acid group on the end of the compound structure where the reduction reaction occurs. As stated previously, after the reduction reaction oc curs, the site reactivates via rehydration or oxidation. Over time, however, the catalyst activity decreases, requiring replacement, washing/cleaning, rejuvenation, or regeneration of the catalyst. Catalyst designs and formulations are generally proprietar y. Both the catalyst material and configuration determine the properties of the catalyst.

Originally, SCR catalysts were precious metals such as platinum (Pt). In the late 1970s, Japanese researchers used base metals consisting of vanadium (V), titanium (T i), and tungsten (W), which significantly reduced catalyst cost. In the 1980s, metal oxides such as titanium oxide (TiO 2 ), zirconium oxide (ZrO 2 ), van a dium pentoxide (V 2 O 5 ), and silicon oxide (SiO 2 ) were employed to broaden the reaction temperature range. Zeolites, crystal line alumina silicates, were also introduced for high temperature (675

– 1000°F; 360

– 540°C) applications; however, zeolitesChapter 2

– Selective Catalytic Reduction

2

- 14 tended to be cost prohibitive

. From 1980 to 2008 , the cost of catalyst has dropped from approximately $34,000/m 3 to a range of $5,000 to $6,000/m 3 (costs are in 2011 $) [ 7 b ]

. 5 This reference also reported that catalyst prices remained in the approximate range of $5,000 to $6,000/m 3 through 2012.
Improvements to the catalyst formulations over time have decrease d unwanted side reactions such as sulfur oxide conversions ( SO 2 to SO 3 ) and increase d the resistance to flue gas poisons , and newer catalysts can oxidize metallic mercury (Hg) into ionic forms (for easy removal downstream in wet scrubbers and wet electrosta tic precipitators [ ESPs ] ) [ 4 0 ]

. Improved

c atalyst designs have also increase d catalyst activity, surface ar ea per unit volume, and the
tem perature range for the reduction reaction. As a consequence, there is a corresponding d e crease in the required catalyst volumes and an increase in the catalyst operating life. For coal

fired boiler applications, SCR catalyst vendors typically guarantee the catalyst for an operating life ranging from 8,000 to 24,000 hours [ 1 ]. Applications usi ng oil and natural gas have a longer operating life, typically greater than 32,000 hours [ 4 1 ]. In addition, operating expe rience indicates that actual catalyst deactivation rates are lower than the design specifications [ 33 ]. The latest demands on catalyst technology for both higher and lower sulfur coal

- fired boilers include design NO x removal of 90 % ; control of residual NH 3 to 2 ppm (i.e., ammonia slip) ; guarantees for SO 2 oxidation to less than 1 %

, and in many cases

, to less than 0.5 % ; being able to with stand washing/cleaning and regeneration procedures ; and guarantees for mercury oxidation [ 37 ]

. Catalyst formulations include single component, multi

- component, or active phase with a support structure. Most catalyst formulations cont ain additional compound s or sup ports to give thermal and structural stability or to increase surface area [ 4 2 ]. Catalyst configurations are generally ceramic honeycomb and pleated metal plate (monolith) designs in a fixed

- bed reac tor, which provide high surface area to volume ratio. Pellet catalyst s in fluidized beds are also available. Pellets have greater surface area than honeycombs or pleated plates but are more susceptible to plugging. This lim its the use of pellets to clean

- burning fuels such as natural gas.
Catalyst elem ents placed in a frame form a catalyst module. The modules stack together in multiple layers to create a reactor bed of the total required catalyst volume. A typical module is 3.3 ft × 6.6 ft in area (1 m × 2 m) and 3.3 ft (1 m) in height. A crane hoists t he large catalyst modules into the reactor from either the interior or exterior of the reactor, depending on the reactor design.
Catalysts greatly accelerate the NO x reduction reaction rate, but some catalysts have more favorable properties for a given app lication. Performance requirements that drive the choice of catalyst include reaction temperature range, flue gas flow rate, fuel source, catalyst activity and selectivity, SO 2 oxidation

, and catalyst operating life. In addition, the design must consider t he cost of the catalyst, including disposal costs [ 33 ]. In the past, the initial charge of catalyst costs
account ed for 20% or more of the capital costs for a n SCR system [ 33 ] , however, as catalyst unit cost has declined over time, th is catalyst cost is a smaller percentage of the capital costs [ 7 ]

. The catalyst layers may be washed/cleaned, rejuvenated, or regenerated to extend the catalyst life as catalyst activity declines , or they may be replaced

. Generally, less than one layer

5 An earlier reference shows that f rom 1980 t o 2006 , the cost of catalyst dropped by 75% from a pproximately $16,000/m 3 to less than $4,000/m 3 [37]. These costs are for the cost year reported and are not adjusted for escalation to current year.Chapter 2

– Selective Catalytic Reduction

2

- 15 of catalyst is replaced per year for fixed bed designs. Most SCR manufacturers offer a disposal service

, in which either t he catalyst is reacti vated (i.e., rejuvenated or regenerated) for reuse or its components are recycled for other uses [ 4 2 ]. If the catalyst cannot be recycled or reused, the facility operator must dispose of the spent catalyst in an approved landfill. In the United States, most catalyst formulations are not considered hazardous waste [ 4 2 ].
Catalyst cleaning typically means the removal of physic al restrictions t o the catalyst (i.e., blinding layers and large particle ash [ LPA ] )

, rejuvenation means the removal of poisons without replenishing catalytically active compounds in the catalyst, and regeneration typically means the removal of poisons and the restoration of catalytic activity by restoring catalyst active ingredients [ 4 3 ]

. These activities may occur online/in

- situ or offline

, and may occur onsite or offsite [ 4 3 ]

. The catalyst layers may be removed and transported to the cleaning, rejuvenation, or regenerati on site

. Online catalyst cleaning with soot blowers or sonic horns is conducted on a regular basis to remove ash or particles ( soot blowers and sonic horns are discussed below ) [ 4 4 ]

. Water

- based cleaning can also be conducted to remove physical materials that plug or blind the catalyst [ 4 5 ]

. Use of rejuvenated and regenerated catalyst has increased since the late 1990s

, and for some applications

, can be considered equivalent to new ca talyst [ 4 6 ]

. For rejuvenation processes, the focus is on removal of blinding materials and catalyst poisons [ 43, 47 ]

. Rejuvenation processes may cause the catalyst to lose structural integrity and mechanical strength [ 4 3 ]

. Today, regeneration processes foc us on increasing the l ongevity of the catalyst by maintaining its mechanical strength (both compressive and bonding) and improving its activity or performance [ 43, 47 ]

. The catalyst layer may be washed in a series of baths that remov e ash and particulate,

remove poisons, and add chemicals or various metals used in the original catalyst manufacturing [ 45, 47 ]

. Moisture must also be removed from the catalyst because it reduces the strength of the cataly st [ 4 7 ]

. Because regeneration may cause catalysts to lose mechanical strength

, recalcination ma y also be conducted to ensure catalyst impregnation and that catalyst mechanical streng th is regained [ 43, 45, 47 ]

. Improvements to the catalyst activity relative to certain reaction s can be made using a regeneration process (e.g., impregnating other chemicals in the catalyst to reduce the conversion of SO 2 to SO 3 , or to increase the oxidation of Hg over the conversion rate of the original catalyst while maintaining the same catalyst a ctivity for NO x ) [ 4 7 ]

. Damage to the catalyst can occur during operation of the SCR or during transport of the catalyst for rework, so a thorough inspection of the modules is conducted along with replacement

or repair of any damaged elements [ 4 4 ]

. Benefits for regenerated catalysts include the following: lower regeneration cost per layer as compared to new replacement cost [1]; no catalyst cost [1]; disposal cost savings [1]; full restoration of original catalytic activity [McMahon Feb 2006 for Platts ]; sam e deactivation rate as a new catalyst in the same SCR installation, i.e., comparable equipment life as new catalyst [McMahon Feb 2006 for Platts ]; lower SO 2 oxidation for some catalysts (i.e., SO 2 /SO 3 conversions is no higher than the new catalyst guarante e) [50]; removal of physical restrictions, including fly ash plugging, large particle ash, and blinding layers [43]; and no physical damage to the catalyst and no loss of structural integrity [43].

Use of r egenerated catalyst s reduces catalyst replacement co st and minimizes the need to dispose of spent catalyst [ 4 8 ]

. It is estimated that a typical 500

- MW coal

- fired power plant willChapter 2

– Selective Catalytic Reduction

2

- 16 spend approximately $2,000,000 6 on a single layer of new catalyst [ 4 8 ]

. The cost for regenerated catalyst for this same facility would be approximately $1,000,000 for a single layer of catalyst [ 4 8 ]

. Disposal costs when replacing a spent catalyst could be $50,00 0

– $200,000 per layer, and these costs are avo ided with regenerated catalysts [ 4 8 ]

. Regenerated catalyst typically costs 40 percent less than new catalyst [ 48

- 50 ]

. 2.2.2 SCR Performance Parameters

The rate of the reduction reaction determines the amount of NO x removed from the flue gas. The major design and operational factors that affect the NO x removal performance o f SCR are similar to those presented in Chapter 1

, SNCR. The factors discussed previously for SNCR include the following:

 Reaction temperature range;

 Residence time available in the optimum temperature range;

 Degree of mixing between the injected reagent a nd the combustion gases;

 Molar ratio of injected reagent to inlet NO x ;

 Inlet NO x concentration level; and

 Ammonia slip.

The majority of the discussion regarding SNCR design and operational factors is valid for the SCR process

, except for small variations due to the use of a catalyst and the reaction chamber being separate from the combustion unit. Additional design and o pera tional factors to consider that are specific to the SCR process include the following :

 Catalyst activity ;

 Catalyst selectivity ;

 Pressure drop across the catalyst ;

 Ash management (i.e., mitigating large particle ash (LPA) impacts on the catalyst) and dust loading ;

 Catalyst pitch ;

 SO 2 and SO 3 concentrations in gas stream;

 Catalyst deactivation ; and

 Catalyst management

. The major differences between SNCR and SCR are discussed below.
Temperature

The NO x reduction reaction is effective only within a given temperature range. The use of a catalyst in the SCR process lowers the temperature range required to maximize the NO x re duction reaction. At temperatures below the specified range, the reaction kinetics decrease and ammonia passes through the boiler (ammonia slip) , but there is little effect on nitrous oxide

6 Cost year not available; data are from 2008 article [Re ference 48 ].Chapter 2

– Selective Catalytic Reduction

2

- 17 (N 2 O) formation . At temperatures above the specified range, N 2 O fo rm ation increases and catalyst sintering and deactivation occurs , but little ammonia slip occurs

. In an SCR system, the optimum temperature depends on both the type of catalyst used in the process and the flue gas composition. For the majority of commercial cata lysts (metal oxides ), the operating temperatures for the SCR process range from 480 °F to 800°F (250

– 430 °C) [ 4 6 ]. Figure 2.2 is a graph of the NO x removal efficiency as a function of temperature for a typical metal oxide catalyst [ 4 6 ]. The fi gure shows that the rate of NO x removal increases with temperature up to a ma ximum between 700°F and 750°F (3 70

– 400°C). As the temperature increases above 750°F (400°C) , the reaction rate and resulting NO x re moval efficiency begin to decrease.

Figure 2.2 : NO x Removal versus Temperature [ 4 6 ]

As flue gas temperature approaches the optimum, the reaction rate increases and less catalyst volume achieves the same NO x removal efficiency. Figure 2.3 shows the change in the required catalyst volume versus temperature [ 51 ]. There is approximately a 40% decrease in the required catalyst volume as flue gas temperature increases from 600 ° F (320 ° C) to the optimum
rang e, 700

– 7 50°F (370

– 400 °C). This decrease in catalyst volume also results in a significant decrease in capital cost for the SCR system. Less catalyst also results in a decrease in annual operation and maintenance costs. For example, the system pressure drop would be lower, which would reduce the additional electricity needed to run the ID fan. The net effect on catalyst replacement costs is uncertain; although the volume of catalyst replaced would be smaller, deactivation may occur more frequently since the quantity of materials in the emission stream responsible for plugging and poisoning would not be reduced.Chapter 2

– Selective Catalytic Reduction

2

- 18 The relationships between flue gas te mperature, catalyst volume, and NO x removal are complicated functions of the catalyst formulation and configuration. The p hysical and chemical properties of each catalyst are optimized for different operating conditions . For a given catalyst formulation, the required catalyst volume or temperature range can even change from one manufacturer of the catalyst to another. The refo re, the selection of the catalyst is c ritical to the operation and performance of the SCR system.

Figure 2

. 3: Change in Catalyst Volume vs. Temperature [ 51 ]

Because the optimum temperature window of the SCR process is lower than that of SNCR, the reagent injection into a reactor chamber occurs downstream of the combustion unit, rather than inside the combustion unit. As discussed previously, there are several options for the location of the SCR reactor. The flue gas temperature at each of these locations is different. Most designs install the reactor downstream of the economizer and prior to the air preheater, where the flue gas is at the appropriate temperature for m etal oxide

– based catalysts. Reheating of the flue gas may be required for reactors locate d downstream of the air pre heater. Reheating significantly increases SCR operational costs. This continues to be true despite natural gas prices that are relatively lo w on an historical basis. 7

7 A case study of tail

- end SCR indicated that the cost of natural gas for flue gas reheating was about 60% of the variable annual operation and maintenance costs (i.e., sum of natural gas, electricity, and reagent costs) when the unit cost of natural gas was $8/1,000 sft 3 [52]. If the natural gas unit cost were $5/1,000 sft 3 , then the natural gas cost for flue gas reheating would have been more than 40% of the variable annual operation and maintenance
costs.Chapter 2

– Selective Catalytic Reduction

2

- 19 Boiler operation at reduced loads decreases the gas flow rate. At reduced gas flow rates, the economizer outlet gas temperature decreases because boiler heat transfer surfaces absorb more heat from the flue gas. Typical SCR syste ms tolerat e temperature fluctuations of ±200°F (± 93°C) [ 33 ]. At low boiler loads, however, the temperature can decrease below the optimum range. For example, a coal

- fired utility boiler has an economizer exit flue gas temperature of 690°F ( 370 °C) at 100% l oad, but only 570°F (300°C) at 50% load [ 33 ]. For low

- load operations, an economizer bypass can be used to raise the flue gas temperature . An economizer bypass diverts part of the hot flue gas from within the economizer through a bypass duct and mixes it

w ith the relatively cooler flue gas exiting the economizer. An economizer feedwater bypass also raises the flue gas t emperature. The use of an econo mizer bypass results in less energy transfer to the feedwa ter for steam generation; conse quently, there is a small reduction in boiler efficiency. Lower boiler efficiencies require more fuel to be burned to meet the required boiler steam output.

Re sidence Time and Space Velocity
Residence time is the time the reactants are within the reactor. Higher residence tim es generally result in higher NO x removal rates. Temperature also affects the required residence time. The required residence time decreases as the temperature approaches the optimum temperature for the reduction reaction. Residence time is often expressed as space velocity, the inverse of residence time. The space velocity of a reactor is experimentally determined from the measured flue gas flow rate divided by the superficial volume of the catalytic reactor. The NO x removal efficiency increases with decre asing space velocity ( i.e.

, increasing catalyst volume ) for a given flue gas flow rate.
The optimal residence time for an SCR system is a function of the number of active catalyst sites available for the reduction reaction and the gas flow rates within tho se active sites (interstitial flow rate). The “ area velocity ” is a parameter used by SCR vendors that relates the number of sites and the interstitial flow rate to residence time. The area velocity is defined as the space velocity divided by the catalyst p ore surface area (specific surface area). For coal

- fired boilers, typical specific surface areas range from 90 to 3,800 square feet per cubic feet (ft 2 /ft 3 ) (300

– 1,200 square meters per cubic meters [ m 2 /m 3 ] ) [ 33 ]. In creasing the catalyst specific surface area increases the NO x removal for a given flue gas flow rate. This can be accomplished by either increasin g the catalyst volume, which in creases the reactor size, or increasing the pore space of t he catalyst, which generally in creases the catalyst cost.
Degree of Mixing
The reagent must be dispersed and mixed throughout the flue gas to ensure suffi cient contact between the reactants. Mixing is performed by an injection system that injects pressurized gas

- phase ammonia into the flue gas. The injection system controls the spray angle, velocity

, and direction of the injected reagen t. Some systems inject the ammo nia with a carrier fluid

, such as steam or air

, to increase penetration in to the flue gas. Injection systems are application specific. Numeric modeling of the flue gas and reagent flow opti mizes the design of the injection system ( s ee Section 2.2.6).
Mixing of the flue gas and ammonia occurs before entering the SCR reactor. If m ixing is not adequate, the NO x reduction is inefficient. SCR designs must incorporate adequate ductChapter 2

– Selective Catalytic Reduction

2

- 20 length between the ammonia injection and the reactor inlet to allow for mixing. Mixing patterns can be improved by

 Installation of static mixers upstream of the reactor;

 Increas ing the energy imparted to the injected fluids;

 Increasing the number of injectors and/or injection zones; and

 Modify ing the nozzle design to improve the reagent distribution, spray angle, and direction.

Stoichiometric Ratio Factor

The stoichiometric ratio factor (the moles of reagent injected per mole of NO x removed ) defines the quantity of reagent needed to achieve the targeted NO x reduction. Theoretically, based on reaction equations 2.1(a) and (b), one mole of NO can be removed with one mole of ammonia and one mole of NO 2 can be removed with t wo moles of ammonia

. Since NO x is mostly comprised of NO (approximately 95%), the theoretical stoichiometric ratio for NO x is close to 1.05 mole of ammonia per mole of NO x . H ence, SCR systems typically employ a stoichiometric ratio of 1.05 moles of ammonia per mole of NO x [ 33 ]. This assumption of a n almost one

- to

- one linear relationship between the quantity of reagent and the NO x removed is good up to about 85% NO x reduction [ 4 6 ]. A bove 85%, the removal efficiency begins to level off and more than the theoretical amount of ammonia is required for additional NO x re moval because of reaction rate limitations. Because capital and operating costs depend on the quantity of reagent con sumed, the stoichiometric ratio factor is an important design parameter that is determined by the SCR designer.

Inlet NO x Concentration

The concentration of the reactants also affects the reaction rate of the NO x reduction process. In general, higher i nlet NO x inlet concentrations result in higher NO x re moval efficiencies due to reaction kinetics [33] . However, NO x concentrations higher than approxi mately 150 parts per million (ppm) generally do not result in increased performance. Low NO x inlet levels result in decreased NO x removal efficiencies because the reaction rates are slower, particularly in the last layer of catalyst [ 33 ]. The percent removal efficiency achieved is dependent on the inlet NOx concentration, so that SCR that follow other NOx cont rols such as LNB or FGR may achieve an efficiency less than 90%. In general, though, SCR achieves greater removal efficiencies than SNCR on sources with low inlet NO x levels , such as natural gas

– fired boilers.
For a given NO x removal efficiency, higher NO x levels at the SCR inlet require more catalyst volume. For example, to achieve 90% NO x removal requires 10% more catalyst at an inlet NO x level of 1.7 lb/MMBtu versus an inlet NO x level of 0.8 lb/MMBtu [ 5 3 ].
Ammonia Slip
Ammonia slip refers to the excess reagent passing through the reactor. Ammonia in the flue gas causes a number of problems

, which were disc ussed in Chapter 1

, SNCR, includ ing health effects, visibility of the stack effluent, salability of the fly ash, and formation ofChapter 2

– Selective Catalytic Reduction

2

- 21 ammonium sulfates. Li mits on acceptable ammonia slip, imposed by either regulatory limits or design requirements, place constraints on SCR performance.
Ammonia slip does not remain constant as the SCR system operates but increases as the catalyst activity decreases. Properly d esigned SCR systems, which operate close to the theoretical stoichiometry and supply adequate catalyst volume, maintain low ammon ia slip levels, approximately less than 2 ppm [ 1 ]

. While ammonia slip levels in operating permits are typically in the range of 2 to 10 ppm, in actual practice lower slip levels are achieved, and the slip levels approach permitted levels only when the catalyst is near the end of its service life [ 1 ]. Ammonia slip increase s with the NSR ratio. 8 Ammonia slip monitoring instruments are commercially available and are in place and operating at a number of coal

- fired units. Facilities typically install ammonia slip monitors between the SCR and the air heater and may measure at one or several points. These systems monitor ammonia slip an d help the unit maintain slip levels of 2

– 3 ppmv or less. The capital cost for one ammonia slip monitoring instrument is estimated to be $40,000 for a single measurement point and up to $70,000 in capital cost for three measurement points [ 5 4 ]. Another method for quantifying ammonia slip is to determine the ammonia concentration in collected fly ash [ 5 5 ].

Raw materials at some cement kilns contain constituents that r elease ammonia to the kiln gas stream when heated to high temperature. S ome cement plants have ammonia in the kiln exhaust gas without injecting any ammonia into the gas stream. Therefore, the ammonia slip from unreacted ammonia injected for SCR is difficult to differentiate from the natural fluctuations in ammonia in the stack gas. For this reason , it is important to understand the level of raw material derived ammonia emissions when designing an SCR control system for cement kilns [94]

. Catalyst Activity
Catalyst activity is a measure of how much the catalyst accelerates the NO x reduction reaction rate. A higher catalyst activity results in a faster reaction rate and more NO x removal. Catalyst activity is a function of many variables

, including catalyst composition and structure, diffusion rates, mass transfer rates, gas temperature, and g as composition [ 5 6 ]. As the catalyst activity decreases, the NO x reduction reaction rate also decreases. This results in lower NO x removal and higher ammonia slip levels.
The following equation describes the deactivation of the catalyst activity, K, with
t ime, t [ 5 7 ]:

(2.2)

where K o is the original catalyst activity and  is the catalyst opera ting life time constant. Figure 2.4 shows a typical catalyst deactivation curve based on Equation 2.2. As the catalyst activity decreases, the NO x removal efficiency is usually kept constant by injecting more ammonia, thereby increasing the ammonia slip. When the ammonia slip reaches the maximum

8 Lewandowski, D. Design of Thermal Oxidation Systems for Volatiles Organic Compounds

.Chapter 2

– Selective Catalytic Reduction

2

- 22 design or permitted level, the catalyst or a catalyst layer must be cleaned, rejuvenated

, or regenerat ed, or new catalyst must be installed.
Catalyst Reaction Selectivity

SCR favors the NO x reduction reaction over competing reactions if the reactants are at the appropriate temperature and oxygen is present. However, competing reactions still occur

, and the catalyst accelerates these reactions as well. Each catalyst has different chemical reaction selectivity properties. In general, catalysts promote the formation of two undesirable compounds, sulfur trioxide (SO 3 ) and nitrous oxide (N 2 O). SO 3 is for med by the oxidation of SO 2 to SO 3 , which occurs both during combustion of sulfur

- containing fuel and over the catalyst. In cement plant applications, the fuel sulfur is incorporated into the clinker, however, pyritic sulfur contained in some raw material s is oxidized and released into the kiln gas stream [94]

. Sulfur oxides (SO x ) are regulated under the 1990 Clean Air Act. SO 3 reacts with ammonia in the flue gas to form ammonia sulfates. Ammonium sulfur salts depo sit on the catalyst and on down stream equi pment such as the air preheaters

. SO 3 and a mmonium sulfate formation is primarily a concern for higher sulfur coals. Additional costs for air preheater modifications are included in the cost analysis when the sulfur content of fuel is greater than 2 percen t or the SO 2 content of the fuel is greater than or equal to 3 lb/mmBtu. Increasing the number of catalyst layers, while increasing NO x removal efficiency, can also lead to an increase in the conversion of SO 2 to SO 3

. 9 In addition, SO 3 emissions cause “blue plume” from the unit and have become an emissions concern. Newer catalysts , however, have been developed that limit the formation of SO 3 [ 37 ] , and these catalysts are now commonly used in SCR installations in the U.S . N 2 O is both an ozone depleter and a greenhouse gas. N 2 O has a global warming potential (GWP) of 298 as compared to carbon dioxide (CO 2 ). 10

9 U.S. EPA. Federal Implementation Plan (FIP) proposal for the Four Corners Power Plant. 77 FR 51620 (August 24, 2012). Correspondence between U.S. EPA and Hitachi Power Systems America.
10 It should be noted that EPA issued a final rule on November 29, 2013 that changed the GWP of N 2 O to 298 (from 310) as part of a notice of data availability concerning the Mandatory Greenhouse Gas Rule. The November 29, 2013 notice can be found in the Federal Register a t http://www.gpo.gov/fdsys/pkg/FR

- 2013

- 11

- 29/pdf/2013

- 27996.pdf

.Chapter 2

– Selective Catalytic Reduction

2

- 23 Figure 2.4: Typical Catalyst Deactivation per Equation 2.2 with K O = 24.12;  = 55,000

Pressure Loss
The flue gas pressure decrea ses as the flue gas flows across the catalyst. The decrease in pressure is a function of the length of the catalyst and the catalyst configuration. Deposition of fly ash and other particulates on the catalyst over time increases this pressure drop across t he catalyst. The flue gas pressure can be increased by installing new draft fans or by upgrading existing fans. To minimize the pressure loss across the catalyst, the SCR reactor ductwork can be expanded and flow rectifiers and turning vanes can be install ed. Pressure loss is of greater concern in turbine applications, which rely on air flow rather than heat transfer to generate
power.

Pressure loss may also be an issue in cement kiln SCR applications as the dust loading in cement kiln s can be in excess of 80 g/Nm 3 [94]. In addition, cement dust can be sticky

, creat ing residue buildups that can be difficult to remove , and generally result in higher pressure loss than observed in other industrial operations

. As such, it may require addi tional catalyst cleaning and catalysts with larger pitch [94]

. Ash Management and Dust Loading

A sh and dust loading in the SCR catalyst may occur from both fuel combustion (e.g. coal) and raw materials ( e.g., cement raw materials) from the process. The am ount of ash or dust loading varies depending on fuel used, source type, and raw material feeds to the process. These particulates can mask and block the catalyst surface and inhibit the SCR NOx reductionChapter 2

– Selective Catalytic Reduction

2

- 24 efficiency. One source indicates dust loading of 6 .7 gr/dscf for a coal

- fired boiler (PRB), dust loading of 4.2 to 15.5 gr/dscf for a typical long wet cement kiln, and 20.1 to 40.2 gr/dscf for a typical PH/PC cement system [111]. One cement industry preheater kiln in Europe has a dust loading of 35.0 gr/ dscf (80 g/m 3 ) [94].
Catalyst Pitch
Catalyst pitch is a term used with honeycomb and metal plate catalyst

, and it affects the flue gas velocity in interstitial spaces [ 33 ]. As shown in Figure 2.5

, pitch ( represented as p ) is the width of the catalyst cell plus the cell wall thickness

, a . For a given flow rate, wider pitch will result in lower interstitial gas velocities. Appropriate catalyst pitch is important to en sure that ash will not deposit and bridge over catalyst cells and pores. Plugging of the cata lyst reduces the effective surface area by decreasing the number of active sites available for the NO x reduction reaction.
For high

- dust applications, such as c ement kilns and coal

- fired boiler s , catalysts with larger openings or pitch should be used to re duce the potential for catalyst plugging

. However,
t he wider pitch reduces the surface area per unit volume, and may necessitate additional catalyst volume to meet emissions limits for NO x and ammonia slip.

Figure 2.5: Pitch for a Honeycomb Catalyst Configuration

SO 2 and SO 3 Concentrations in Gas Streams

Sulfur in the fuel forms SO 2 during combustion in the boiler, i.e., boiler conversion rate

, and some of the SO 2 is further oxidized to SO 3 within the boiler

. In a coal

- fired boiler, approximately 0.5 to 1.5 percent of SO 2 is converted to SO 3 at typical combustion temperatures, however, as the temperature drops more SO 3 can be formed. 11 Additional SO 3 is formed when SO 2 passes through the SCR catalyst (i.e., SCR conversion rate)

. SCR catalysts can be designed for ranges of conversion such as 0.2 to 0.8 percent per catalyst layer ( for 3 layers of catalyst, this would result in conversion of 0.6 to 2.4 percent). SO 3 will combine with water to form sulfuric acid (H 2 SO 4 ) or sulfuric acid mist. Concentrat ions of SO 3 and H 2 SO 4 of 6 to 10 ppm can cause a visible plume, or a blue plume. [Moretti August 2006]

11 Another source indicates that the amount of SO 3 is approximately 0 to 0.2 percent of the total SOx from the combustion process. [Ehrnschwender Aug 2008]Chapter 2

– Selective Catalytic Reduction

2

- 25 Catalyst Deactivation
Catalysts lose their activity over time for various reasons. The primary mechanisms for catalyst deactivation and surface area loss are discussed below.
Poisoning

– Certain fuel constituents that are released during combustion act as catalyst poisons. Catalyst poisons include calcium oxide

, magnesium oxide, potassium, so dium, arsenic, chlorine, fluorine, and lead. These constituen ts deactivate the catalyst by diffusing into active
pore sites and occupying them. Catalyst poisoning repr e sents the main cause of catalyst deactivation. The water soluble alkali compounds are known to react with active catalyst sites rendering them inert. The form of the alkali metals is important as water soluble alkali metals have been found to be more detrimental to catalyst activity.
Arsenic is a concern for boilers that reinject fly ash or burn coals that have moderate to high arsenic content with a free lime (CaO) concentration in the ash of less than 2 percent. Some facilities have found it useful to inject limestone into the furnace. The calcium oxide reacts with arsenic to form calcium arsenate, which affectively removes the arsenic from the gas s tream thereby preventing it from binding to catalyst and rendering them inactive [96]. The activity of poisoned catalysts may be restored by rejuvenation or regeneration.
In cement kilns, both the raw materials and the fuels can contain compounds that pois on catalysts, including inorganic compounds of potassium, sodium, and chlorine. Alkali metal compounds in the raw materials have relatively low melting points and upon reaching the combustion zone of the kiln will readily volatize. Although these compounds are also found in coal

- fired boilers, they may be present in higher concentrations in cement kiln exhausts [94]. Catalyst poisoning by phosphorus, chromium, and lead compounds is believed to be a lesser concern for cement kilns than for coal

- fired boilers [94]. These compounds are expected to be present in lower concentrations in cement kiln exhaust gases than is typical of coal

- fired boilers [24, 94]. Arsenic poisoning is not expected to be an issue for cement plants because the high concentration of CaO should react with any arsenic in the exhaust gas before it reaches the SCR catalyst [96, 111]. However, high CaO levels combined with high sulfur concentrations can result in the formation of CaSO 4 in cement kilns, which can reduce catalyst activity throug h masking [94].
Thermal Sintering

– High flue gas temperatures within the SCR reactor cause sintering, a permanent loss of catalyst activity due to a change in the pore structure of the catalyst. Thermal sintering can occur at temperatures as low as 450°F ( 230 °C). The amount of thermal sintering depends on the composition and structure of the catalyst. Newer catalyst materials are less susceptible to thermal sintering

, which increases their operating life.
Blinding, Plugging , and Fouling

– Ammonia

- sulfur salts, fly ash, and other particulate matter in the flue gas cause blinding, plugging

, or fouling of the catalyst. The particulate matter de posits on the surface and in the active pore sites of the catalyst. This results in a decrease of t he number of sites available for NO x reduction and an increase in flue gas pressure loss across the catalyst.

Erosion

– Impingement of particulate matter and high interstitial gas velocities erode the catalyst material. Catalysts with hardened leading edge s or increased structural strength are lessChapter 2

– Selective Catalytic Reduction

2

- 26 susceptible to erosion. However, i ncreasing catal yst strength through hardening reduces the number of active pore sites.
Aging

– Catalyst aging is a change in the physical and chemical properties of the catalyst pores that occurs over time.

A number of measures can be taken to decrease the rate of deactivation and deterioration of the catalyst. These measures are discussed below

. Catalyst Formulation

– Each catalyst formulation has diffe rent physical and chemical
prop erties. Catalyst formulations with the following properties will have decreased deactivation :

 Increased activity per unit volume ;

 Greater thermal resistance ;

 Chemical and physical resistance to poisons ;

 Wider thermal operating range ;

 Greater structural strength and hardened leading edges ; and

 Lower interstitial velocities (i.e., wider catalyst pitch)

. To obtain the optim al catalyst formulation and SCR design for an application, the catalyst supplier and SCR vendor should be informed of the fuel and raw material constituents , such as sulfur, chlorine, fluorine, alkali metals, lead, arsenic , and trace metals. The SCR must be designed to operate for the full range of values for the constituent concentrations, therefore, the ranges of these data should be provided to the catalyst supplier. These constituents can be determined by chemical analyses. The associated analytical data can then be used to modify the catalyst composition

, determine the appropriate catalyst volume , predict catalyst regeneration

a nd replacement schedules, and design the SCR reactor components.

Soot Blowers and Sonic Horns

– Deposits on the surface of the catalyst can be dislodged by soot blowers

, which are generally installed between each catalyst layer and operated on a periodic b asis, such as once a week

. For some higher dust SCR systems, e.g., in the cement industry, more frequent and even continuous catalyst cleaning may be needed. A sonic horn is an other option that preven ts accumulation of ash deposits on the catalyst surface

. A sonic horn may operate at a typical frequency of 10 seconds every 10 minutes [ 37 ]

. Turning Vanes and Rectifier Grids

– Particulate matter can be removed from the flue gas by gas

- flow turning vanes and flow rectifier grids near the front of the catalyst layer. Particles impact the surface of the vanes or grid and fall out of the flue gas stream. In addition to removing par ticles, turning vanes and flow rectifier grids decrease the linear velocity of the flue gas and align its vector with the flow path of the catalyst.
Catalyst Management Plan
Catalyst deactivation is an inherent part of the SCR process. As the catalyst acti vity decreases with time, the NO x reduction reaction rate decreases and ammonia slip increases. When the ammonia slip level reaches the design limit, the catalyst must be replaced with regenerated catalyst or new catalyst must be added. The catalyst life i s the time the catalyst activity for a given catalyst volume (layer volume) maintains ammonia slip below the designChapter 2

– Selective Catalytic Reduction

2

- 27 limit. Currently, vendor

- guaran teed life for a catalyst layer in coal

- fired applications is typically three years [ 51 ] , and actual catalyst layer lifetimes in such applications are often in the 5 to 7 year range, depending on the condition of untreated flue gas [ 33 ]. Gas

- and oil

- fired applications experience even longer catalyst layer lifetimes.
A catalyst management plan (CMP) , as described in a “saw

- toothed” graphic shown in Figure 2.6, schedules periodic replacement of catalyst to maintain ammonia slip limits (this CMP has a maximum ammonia slip design value of 2 ppm [ 5 8 ]). In the past, CMP descriptions focused mainly on the catalyst replac ement schedules ; however, today, a CMP is a comprehensive catalyst management strategy that incorporates both SCR equipment management and catalyst management , along with attention to changes in regulatory requirements

. This more comprehensive approach is needed with the move to year

- round operation of SCRs for facilities that operate continuously

. While operation of SCR for compliance with the NO x SIP Call (1998) requirements typically called for ozone season operation only, year round operation is necess ary to comply with more recent regulations (e.g., The Clean Air Interstate Rule (CAIR), 1999 Regional Haze Rule, an Ozone Transport Commission ( OTC ) initiative, and state rules such as the North Carolina ( NC ) Clean Smokestacks Rule that took effect in 2009 and the Texas requirements for the Houston area ), to generate NO x credits , or to comply with settlement agreements with the U.S. EPA and Department of Justice

. Continuous, ongoing collection and documentation of d ata on plant loading and cycling , fuel demands and variation, and ongoing NO x performance and SO 2 conversion

, which can then be compar ed to cata lyst activity data

, is conducted to create the plant operating history

. Some companies have developed computer software that collects these data and o ptimizes the costs for catalyst replacement options

. In general, an annual SCR system inspection is conducted on the catalys t , the reactor

, and the complete NH 3 storage and injection system

. Inspection of the catalyst includes a physical inspection along w ith catalyst sampling and analysis on a bench

- scale reactor for activity, SO 2 to SO 3 conversion rate, and pressure drop for each catalyst layer

. Annual ammonia injection grid ( AIG ) tuning and optimization is also conducted to ensure uniform flow rate/veloc ity and uniform NH 3 /NO x molar distribution . Poor distribution of the NH 3 /NO x decreases the NO x reduction and increases the NH 3 slip [ 5 9 ]

. I n situ
measurement s of the catalyst activity have been developed, where NO analyzers installed before and after the catalyst layer and a small supplemental ammonia controller allow increases in the NH 3 /NO x ratio and measurement of inlet and outlet NO x samples, contained to a small area of the catalyst. In situ catalyst activity measurements may be important for year

- round operation of SCR units [ 60 ]

.Chapter 2

– Selective Catalytic Reduction

2

- 28 Figure 2.6: Typical Catalyst Management Plan [ 6 1 ]

Most CMPs call for the SCR reactor design to provide two or more layers filled with catalyst and one or more empty or spare catalyst layers (often called “2:1” design) . When the initial catalyst layers deactivate to the point where ammonia slip reaches the maximum design value, the facility typically adds catalyst to the empty layer. Catalyst addition is managed so th at the total catalyst activity of all the layers (the two or three older catalyst layers plus the new catalyst layers) is sufficient to meet the ammonia slip requirement for a relatively long period of time. As the catalyst continues to deactivate, ammonia slip begins to rise again. When ammonia slip again reaches the maximum design value, one of the older catalyst layers is removed and replaced. The catalyst analysis data identifies which layer should be replaced. With advances in catalyst regeneration, pa rt of a comprehensive CMP is determining whether the catalyst can be regenerated or whether new catalyst must be used. Before a regeneration process is planned, the process should be prequalified on a catalyst sample. If additional catalyst capabilities ar e needed, review of recent catalyst technology advances for newer catalysts that achieve mercury reductions, lower SO 2 conversion rates, and lower load and temperature operation is advised, although some regeneration processes may offer improvements with t hese catalyst capabilities as well. Typically, t he addition and replacement of catalyst layers is coordinated with plant outage periods if at all possible , and outage frequency should be considered in conjunction with the risk considerations for replacing sooner rather than later [ 5 9 ]

. There would likely be additional costsChapter 2

– Selective Catalytic Reduction

2

- 29 or impacts (e.g., due to lost generation or production ) if a facility is unable to coordinate with planned unit outages.
In the past, c atalyst cost was a significant portion of the annual cost of operating an SCR system. Under the latest operating approaches that involve us ing a CMP , only a fraction of the total catalyst inven tory, rather than the entire volume, is replaced at any one time. This distrib utes the catalyst replace ment costs more evenly over the lifetime of the system, and use of regenerated catalyst may also reduce the overall annual costs [ 5 9 ]

. For applications with higher dust loading, such as the dust loading typical for cement kilns, the catalyst management plan may include more frequent catalyst replacement and regeneration schedules than would be typical for low

- dust applications

. 2.2.3 SCR System Configurations
Electric utility and large industrial boiler applications implement several different SCR system configurations, including high

- dust, low

- dust, and tail

- end arrangements

. In a 1997 report, the SCR configurations were reported as 88 % high

- dust SCR, 6 % low

- dust, and 6 % tail

end [ 6 2 ]

. 12 More recently for the U.S, it was report ed that most SCR configurations are high dust, only one facility has a low

- dust SCR, and no tail

- end SCR operate in the U. S. [ 5 2 ]. High

dust is generally considered the most economical and straightforward design provided sufficient space is available to co nstruct the SCR cl o se to the economizer and air pre

- heater. Boiler units with space constraints must consider low

- dust and tail

- end SCR designs. SCR configurations for gas turbine applications depend on the type of engine cycle, such as combined

- cycle or s imple cycle. The various configurations for boilers and gas

- fired turbines are discussed below. In addition, there are tw o different SCR reactor designs, full SCR and in

- duct SCR

, which are also discussed.
High

- Dust SCR
Figure 2.7 shows a high

- dust SCR sys tem for coal

- fired boiler applications. The SCR reactor location is downstream of the economizer and upstream of the air heater and particulate control devices. The flue gas temperature in this location is usually within the optim al temperature window for NO x reduction reactions using metal oxide catalysts. In this configuration, however, the flue gas contains particulates when it enters the SCR reactor.

Coal

- fired boilers generally use a vertical SCR reactor , where the flue gas flows down ward through the c atalyst. The reactor generally contains multiple layers of catalyst. The volume of catalyst required varies with each installation, as discussed previously. Soot blower s or sonic horns are installed to remove particulates from the catalyst surfaces. For de signs that use a honeycomb catalyst, the cata lyst pitch is typically about 7

– 9 mm (compared with 3

– 4 mm for
gas

- fired boilers) to allow easy passage of ash particles without deposition and ease of cleaning with soot blower s or sonic horns. To obtain uniform gas flow and remove particulates, high

- dust SCR designs usually include turning vanes and a flow

- rectifying grid in the ductwork prior to the
reactor. High

- dust SCR typically require 3 or 4 layers of catalyst [ 5 2 ].

12 In a 2006 report, one utility/vendor reported that of their 24 SCRs, 71% were high

- dust, 4% were low

- dust, and 25% were tail

- end [ 63 ]. These data are from a single vendor; the data above in the text represent multiple vendors.Chapter 2

– Selective Catalytic Reduction

2

- 30 A hopper at the bottom of the SCR reactor collects ash and particulates separated from the flue gas stream. The hopper outlet connects to the plant fly ash handling system for periodic removal of the accumulated ash. Flue gas exits the reactor via an opening at the top of the hopper and i s directed to the air heater inlet. Some designs eliminate the need for hoppers by keeping flue gas velocities high enough in these areas that fly ash remains entrained in the flue gas.
Natural gas

– and distillate oil

– fired boilers generate flue gas that is relatively free of dust and SO 2 (for low

- sulfur oil). Consequently, SCR systems for these boilers place the reactor upstream of the air heater, in the high

- dust SCR configuration.

Figure 2.7: High

- Dust SCR Arrangement [ 4 2 ]

Low

- Dust SCR

Coal

- fired units with an ESP located upstream of the air heater (hot

- side ESP) typically use a low

- dust SCR configuration. Figu re 2.8 shows a low dust configu ration, which locates the SCR reactor downstream of the ESP. In this location, the flue gas is relatively dust fre e. The ash removed by the ESP typically contains arsenic, alkali metals, and other constituents that are detrimental to catalyst performance and life.

A low

- dust SCR system increases catalyst life by mitigating concentrations of particulates and catalyst p oisons in the SCR reactor. In addition, low

- dust SCR configurations do not need ash hoppers. For designs employing honeycomb catalyst, the catalyst pitch ca n be reduced to approximately 4

– 7 mm, resulting in lower catalyst volume. Low

- dust SCR typicallyChapter 2

– Selective Catalytic Reduction

2

- 31 req uires only 2 layers of catalyst [ 5 2 ]. Longer catalyst life, lower catalyst volume, and the elimination of the ash hopper mean lower costs for low

- dust SCR compared to high

- dust configuration s. The only disadvantage of low

- dust SCR is the temperature drop o f the flue gas as it flows through the ESP. Flue gas temperatures generally do not decrease to the point where reheating is required. However, an increase in the size of the existing economizer bypass duct may be required to maintain the flue gas temperatu re within the optimal range.

In the low

- dust SCR installed at a U.S. cement kiln in 2013, the gas stream passes through a hot electrostatic precipitator to remove the majority of the dust prior to entering the SCR. The gas stream exiting the SCR may then pass through a second, more efficient particulate control device to remove the remaining particulate to acceptable emissions rates.

Figure 2.8: Low

- Dust SCR Arrangement [ 4 2 ]

Tail

- E nd SCR

The tail

- end SCR configuration places the SCR reactor downstream of all air pollution control equipment installed on a unit

. Figure 2.9 depicts a tail

- end system for a plant with a particulate control device and a wet FGD system. The air pollution control equipment removes most flue gas constituents detrimental to SCR cat alyst s be fore the flue gas enters the SCR
reactor. The tail

- end SCR configuration is often a technically feasible alternative for situations where the high

- dust SCR configuration is impractical [100].

B ecause the flue gas temperature a t the tail end is b elow the range required for the NH 3 / NO x reaction, the flue gas must be reheated. Tail

- end SCR systems use oil

- or natural gas

– fired duct burners or steam coil gas heaters for reheating. Some of the energy used to reheat the gas is recovered in a recuperati ng gas

- t o

- gas heater. Some use catalysts specially designed forChapter 2

– Selective Catalytic Reduction

2

- 32 temperatures between 300

– 550 o F and low pressure drops, which lowers the costs flue gas reheating [103, 106, 107].

A tail

- end system may have higher capital and operating cost s than the other SCR systems because of the additional equipment and operational costs associated with flue gas reheating and heat recovery. However, these costs are in part offset by reductions in catalyst
costs. . Tail

- end units require less catalyst because they can use catalysts with smaller pitch and higher surface area per unit volume. Tail

- end SCR typically require only 2 layers of catalyst [ 5 2 ] , although some use four half

- layers of catalyst to allow for greater flexibility for catalyst replacement [101] . In addition, because there is less fly ash, catalyst poisons, and SO 2 in the flue gas for tail

- end units, the catalyst lifetime is significantly increased and less expensive catalyst may be used [ 5 2 ]. Some sources have repor ted catalyst lifetimes for tail

- end SCRs to be over 100,000 hours [52, 100, 101]. The tail

- end SCRs may also have longer lifetimes due to the lower operating temperatures and lower levels of dust and SO 3

. Tail

- end SCRs have been used since the late 1980s and were initially used on coal

- fired power plants. They are currently used at a variety of different applications in Europe, Japan, and the U.S., including power plants, incinerators, refinery crackers, cement plants, and ethylene crackers [100, 101]. Th ey have been installed on units burning a wide range of fuels, including fuels of variable composition, such as biomass (including wood waste and chicken litter), hazardous waste, municipal waste, and wastewater sludge [104, 105]. They are often easier and less complex to install than the high

- dust and low

- dust SCR configurations for retrofit situations and can be installed with less disruption to production. The tail

- end SCR configuration has been used in many retrofits of existing power plants in Europe. In some situations, particularly where combustion units have space constraints, the capital cost for retrofitting high

- dust SCRs may be higher than for tail

- end SCR [52]. Modular tail

- end SCR systems are also available that are designed to be installed wi th minimal plant disruption [102].
One other major advantage of the tail

- end SCR configuration is that its preheater enables the SCR to operate independently of the combustion unit. This arrangement enables greater operating flexibility, allowing the comb ustion unit to operate in a wider range of operating loads and fuel types [100]. Because tail

- end units follow the ESP and wet scrubber, the flue gas has cooled and must be reheated to an appropriate temperature for the NO x reaction to occur in the SCR. For tail

- end units, the flue gas is typically sent through a gas

- gas heat exchanger and then to either a natural gas

- fired duct burner or steam coil to heat to the appropriate SCR operating temperature. Most tail

- end SCR in Europe use steam coil reheating , which has advantages over a duct burner such as lower operating cost, no increase in flue gas flow rate from combustion byproducts, and no moisture condensation on the SCR catalyst. 13

13 A case study for a tail

- end SCR achieving 84 percent NO x removal efficiency on a 600 MW boiler burning bituminous coal indicated annual reheating cost for steam coil of $2.5 million/yr and for natural gas burner of $12 million/yr (2008$) (assuming approximately $4/1000 lb steam and $8/1000 sft 3 natural gas) [52]. For comparison

, the annual reheating cost for natural gas burner would be $7.8 million/yr (assuming approximately $5/1000 sft 3 natural gas).Chapter 2

– Selective Catalytic Reduction

2

- 33 Figure 2.9: Tail

- end SCR Arrangement [ 4 2 ]

Gas Turbines

Natural gas

– fired turbine applications frequently use SCR technology for post

combustion NO x control. There are two basic gas turbine configurations: combined cycle (cogeneration cycle) and simple cycle. The majority of SCR systems are installed as combined cycle applications. As shown in Figure 2.10, a typical combined

- cycle SCR design places the reactor chamber within a cavity of the heat recovery steam generator system (HRSG)

, between the superheater and the evaporator . The flue gas temperature in this area is w ithin the operating range for base metal catalysts. Most new HRSG units include a cavity designed to accommodate an SCR reactor. However, older HRSG units may not have sufficient space to house the SCR reactor within the HRSG. In these cases, a low

- tempera ture SCR reactor may be installed after the HRSG. The high temperature SCRs used on simple

- cycle turbines are generally not retrofitted to combined cycle turbines equipped with HRSG due to lack of space between the turbine and the HRSG [95, 96, 97]. Simple

- cycle applications of SCR place the reactor chamber directly at the turbine exhaust, where the flue gas temperature is in the range of 850°F to 1000°F (450

– 540°C). This requires the use of a high

- temperature catalyst such as zeolite [ 4 2 ]

.Chapter 2

– Selective Catalytic Reduction

2

- 34 Figure 2.10: SCR Arrangement for a Combined

- Cycle Gas Turbine [ 4 2 ]

Cement Kilns
SCR systems applied to cement kilns can have “tail

- end”, “low

- dust”, or “high

- dust” configurations. Because of the potential for catalyst plugging, the “high

- dust” configuration on cement k ilns require catalyst cleaning mechanisms. The “low dust” and “tail

- end” configurations avoid the costs of catalyst cleaning systems. Currently, three “high

- dust” SCR systems 14 , three “low

- dust” SCRs 15 , and one “tail

- end” SCR 16 are known. The “high

- dust” SCR s reportedly achieve control efficiencies of approximately 80% with inlet dust loading of up to 100 g/m 3 . The “low

- dust” SCRs are reported to have dust loadings less than 20 mg/m 3 , while the inlet dust loading for the “tail

- end” SCR is reported to be less than 10 g/m 3 [94, 98, 99] SCR Reactor Designs
The reactor design affects the capital and operating costs of the SCR system and the CMP . There are two different types o f SCR reactors : full SCR and in

- duct SCR. Full SCR designs house the catalyst in a sepa rate reactor chamber. The boiler flue gas must be du cted from the economizer outlet to the SCR reactor, then to the air heater inlet. A separate reactor allows a large volume of catalyst to be installed in layers, which increases NO x reduction and catalyst lifetime. It also increases the duct length available for the mixing of reactants before entering the

14 The first “high

- dust” configuration SCR was installed on a preheater cement kiln at the Solnhofer Zementwerkes in Germany in 2001 and operated until 2006. Two other “high

- dust” SCRs have been installed on preheate r cement kilns at the Cementeria di Monselice plant in Bergamo, Italy in 2006 and the Mergelstetten plant in Germany in 2010.
15 “Low

- dust” configuration SCRs have been installed at the Sarche plant in Italy (2007), the Mannersdorf plant in Austria (2012), and the Joppa plant in the USA (2013). The Mannersdorf SCR is installed on a preheater cement kiln, while the Joppa SCR is installed on a long dry kiln. Both plants use an electrostatic precipitator to reduce particulate emissions entering the SCR. The Sar chi SCR is installed on a small Polysius Lepol kiln with no particulate controls, but low dust loading (reportedly less than 15 g/m 3 ).
16 The Rohrdorf plant in Germany installed a “tail

- end” SCR in 2011 on a preheater kiln.Chapter 2

– Selective Catalytic Reduction

2

- 35 reactor chamber. However, a separate reactor requires a large amount of space adjacent the boiler to install the reactor and ductwork. The additional duc twork often necessitates upgrades to the draft fan system.

In

- duct (in line) SCR systems house the reactor within the plant ’ s existing ductwork rather than in a separate reactor chamber. The ductwork is generally enlarged to provide sufficient room for the catalyst. In

- duct systems save on costs for the ductwork, reactor chamber, and induced draft ( ID ) fan. In

- duct designs limit catalyst volume and mixing length ; therefore, they are commonly used in conjunction with other NO x control technologies [ 4 1 ]. Catalyst erosion is generally higher for in

- duct systems. Installation and maintenance of in

- duct systems typically require more boiler outages. Natural gas

– fired boilers, which have low catalyst volumes, frequently employ in

- duct systems. Coal

- fired boile rs frequently employ full SCR reactors but may apply in

- duct SCR reactors where space limitations restrict the installation of a full reactor [ 4 1 ]. Cement kilns have also used full scale SCR reactors, rather than in

- duct SCRs. The SCRs used for cement kiln s have typically consisted of multiple catalyst layers and extensive catalyst cleaning systems. For example, the SCR systems installed at the Solnhofen cement plant in Germany and the Cementeria di Monselice plant in Italy used reactors with six catalyst l ayers, although only three layers were in use at a time [94].

2.2.4 SCR System Primary Equipment
The majority of SCR designs use Thermal De NO x ®, an ammonia

- based NO x reduction system developed and patented by Exxon Research and Engineering Company in 1975

. An SCR system consists of five basic steps :

 Receiv e and stor e the ammonia ( or the urea reactant
, followed by onsite ammonia production ) ;

 Vaporiz e the ammonia and mix it with air;

 Inject the ammonia/air mixture at appropriate locations;

 Mix the ammonia/air with flue gas; and

 Diffuse the reactants into the catalyst and reduc e the NO x . Although the basic steps in an SCR system are similar for all configurations, the system design and equipment specifications are somewhat different. A discussio n of the SCR system design and equipment is given below for an ammonia reagent , high

- dust , full reactor SCR for a 120 MW ( approximately 1,200 MMBtu/hr) coal

- burning utility boiler. These discussions are also pertinent to industrial applications. For exampl e, cement kilns operating in the high

- dust configuration would also require catalyst cleaning equipment [94]. The SCR process steps, related auxiliary equipment, and the potential impacts of SCR operation on existing plant equipment are also discussed. S im plified system flow schematic s are presented in Figure 2.1 and Figure 2.7

, and a list of equipment is presented in Table 2. 3

.Chapter 2

– Selective Catalytic Reduction

2

- 36 Table 2.3 : Major Equipment List for an SCR Application

Item Description/Size SCR reactors (1

– 2) Vertical flow type, 805,000 acfm capacity, 44 ft × 44 ft × 31 ft. high (excluding outlet duct and hoppers), equipped with 9,604 ft 3 of ceramic honeycomb catalyst, insulated casing, soot blowers or sonic horns, hoppers, and hoisting mechanism for catalyst replacement Anhydrous ammonia tank (1 or more) Horizontal tank, 250 psig design pressure, storage tanks 15,000 gal, 34

- ton storage capacity Air compressor (2) Centrifugal type, rated at 3,200 acfm and 30 hp motor Vaporizers (2) Electrical type, rated at 80 kW Mixing chamber Carbon s teel vessel for mixing or air and ammonia Ammonia injection grid Stainless steel construction, piping, valves and nozzles Ammonia supply piping Piping for ammonia unloading and supply, carbon steel pipe: 1.0

- inch diameter, with valves and fittings
Soot blowing steam Steam supply piping for the reactor soot

- piping blowers, 2

- inch diameter pipe with an on

- off control valve and drain and vent valved connections

Air ductwork Ductwork between air blowers, mixing chamber, and ammonia injection grid, carbon s teel, 14

- inch diameter, with two isolation butterfly dampers and expansion joints Flue gas ductwork Ductwork modifications to install the SCR modifications reactors, consisting of insulated duct, static mixers, turning vanes, and expansion joints
Economi zer bypass Ductwork addition to increase flue gas temperature during low loads consisting of insulated duct, flow control dampers, static mixers, turning vanes, expansion joints, and an opening in the boiler casing Ash handling Extension of the existing f ly ash handling modifications system: modifications consisting of twelve slide gate valves, twelve material handling valves, one segregating valve, and ash conveyor piping
Induced draft fans Centrifugal type, 650,000 acfm at 34 inches water gauge (in. w.g .) and 4,000 hp motor
Controls and instrumentation Stand

- alone, microprocessor

- based controls for the SCR system with feedback from the plant controls for the unit load, NO x emissions, etc., including NO x analyzers, air and ammonia flow monitoring devices, ammonia sensing and alarming devices at the tank area, and other miscellaneous instrumentation Electrical supply Electrical wiring, raceway, and conduit to connect the new equipment and controls to the existing plant supply systems Electrical equipment System service transformer OA/FA/

- 60 Hz, 1,000/1,250 kVA (65°C)

Foundations Foundations for the equipment and ductwork/piping, as required
Structural steel Steel for access to and support of the SCR reactors and other equipment, ductwork, and piping

Reagent Production, Storage

, and Vaporization

As discussed previously, one of s everal reagents may be used in an SCR system, including anhydrous ammonia, aqueous ammonia

, or urea

. In the past, reagents h ave typically been purchased and stored before vaporization and use in the SCR

. Ammonia (both anhydrous and aqueous) is the type of reagent most often used in SCR systems. Of about 230 utility boilers for which reagent type was reported in response to a su rvey in 2009, about 80 percent usedChapter 2

– Selective Catalytic Reduction

2

- 37 ammonia, and 20 percent used urea [ 4 ]. Urea reagent is mostly used in SNCR systems [ 6 5 ]

, however, U.S. cement plants typically use 19% aqueous ammonia for SNCR systems and likely would use the same reagent for SCR applic ations. A nother option that some facilities have recently adopted is to produce ammonia onsite from urea feedstock

. The onsite ammonia production system may reduce or eliminate ammonia shipping, handling, and onsite storage

. Load following by the onsite ammonia production sy s tem is extremely important for the proper operation of the SCR.
Several of the pros and cons of each ammonia system are shown in Table 2.4. In general, anhydrous ammonia is the least costly reagent ; however

, plant personnel and community safety, permitting, and other hazard planning concerns associated with its use may make this option less attractive and add to its cost

. Aqueous ammonia is typically higher cost, given the energy required to va porize or decom pose the reagent, although some facilities have chosen this option over anhydrous ammonia to avoid some of the safety and planning concerns for anhydrous ammonia [ 37 ]

. In general, as ammonia consumption increases, onsite urea

- derived ammonia production is the most economical, while for lower consumption rates, aqueous ammonia may be the preferred economic option

. For year

- round operation, onsite urea

- derived ammonia systems become economically competitive with 29 % aqueous ammonia for plants around 800 MW and larger

. For ozone season operation, onsite urea

- derived ammonia systems become competitive with 29 % aqueous ammonia at a plant size of 1,300 MW and larger [ 6 5 ]

. The total cost of an SCR system with an onsite urea

- derived ammonia system is approximately 2 to 5% more than an SCR system based on anhydrous NH 3 [ 10 ]

. Another source reported a capital cost of $24 million for its onsite urea

- derived ammonia system for a 1

, 300 MW unit delivering approximately 7,000 lb/hr NH 3 , with a total capital investment of $175 million for the SCR system (not including the ammonia system) [ 6 6 ]

.

Table 2. 4 : Comparison of Ammonia Delivery Systems [ 6 5 ]

Measure Anhydrous NH 3 Aqueous 19% NH 3 Aqueous 29% NH 3 Urea

- derived NH 3 Risk level Highest safety, hazard, permitting, and regulatory issues Lower safety, hazard, permitting, and regulatory issues Lower safety, hazard, permitting, and regulatory issues Lowest safety, hazard, permitting, regulatory issues

Energy Lowest energy use High energy use Medium energy use Not known
Product deliveries Fewest product deliveries Large number of product deliveries Medium number of product deliveries Not applicable

Capital cost Low capital cost Low capital cost Low capital cost Highest capital cost
Annual cost L owest annual cost High annual cost Medium annual cost Medium annual cost

Aqueous ammonia is typically available as a 19

– 29.4% solution in water. Anhydrous ammonia is nearly 100% pure ammonia and stored as a liquid under pressure. Table 2.2 presents the properties of aqueous (2 9 % ) and anhydrous ammonia.

If f acilities receive anhydrous or aqueous ammonia from offsite , it is received via a tank

truck or rail car and pump ed into one or more storage tanks. Ammonia is typically stored as a liquid in horizontal cylindrical tanks. An a queous ammonia tank is an enclosed tank rated for only slightly elevated pressure, while an anhydrous ammonia tank is a pressure vessel rated for at least 250 pounds per square inch gauge (psig). An anhydrous ammonia tank can be filled to only about 85% of its total volume to allow for a vapor space above the liquid level. The tanks areChapter 2

– Selective Catalytic Reduction

2

- 38 equipped with level and temperature indicators ; a manway, vent, and access ladder ; and other appurtenances. The applicability of heat tracing, insulation, and seismic design criteria are determined based on site

- specific conditions. The tank should be mounted on a concrete pad and surrounded by a spill containment structure such as a dike.
SCR applications on large boilers generally require one to five tanks with volumes
rang ing from 10,000 to 20,000 gallons per tank to maintain sufficient volum e for 1

– 3 weeks of SCR operations. The ammonia storage tank may be sized for 3

– 30 days of storage. The high end of the range would be used in conservative design practice. Alternatively , if ammonia distributors are located nearby and considered reliable, the plant owner might opt for a smaller tank, sized for fewer days of ammonia storage.
Aqueous ammonia is vaporized by pumping it to a vessel where it mixes with hot air. The air from th e dilution air fan is heated in an electric heater or other heat exchanger (e.g., steam). In most aqueous ammonia applications, the ammonia

- air mixture leaves the vaporizer vessel at about 300 ° F (150 ° C) . The vaporization energy required for aqueous ammonia is much greater than that required for anhydrous ammonia because the water in the aqueous ammonia solution also must be vaporized.
If anhydrous ammonia is used, it is fed to the electrical vaporizer by gravity, and the vaporized gas is returned to the sto rage tank vapor space. Vapor is drawn from the vapor space and piped to the ammonia/air mixer. Alternatively, liquid anhydrous ammonia may be pumped to a vaporizer and piped to the ammonia/air mixer.
Ammonia (aqueous or anhydrous), diluted with air at a ra tio of about 20:1 (air: NH 3 ), is transported to the ammonia injection grid. The high proportion of air helps ensure good mixing of air and ammonia and keeps the mixture below the flammable limit.

Onsite Urea

- Derived Ammonia Production

U se of onsite ammonia production systems for feed into SCR reactors has increased

, mostly as a result of safety concerns [ 37 ]

. S everal process types can produce ammonia from urea, including (1) system s that hydrolyze an aqueous urea solution to form ammonia and carbon dioxide (and water); (2) system s that melt solid urea and mix the liquor with steam, where it reacts the melted urea across a catalyst to form ammonia, carbon dioxide, and water vapor; and (3) system s that atomize a n aqueous urea solution in a decomposition chambe r containing a hot air stream or flue gas stream at 800

– 1200 ° F (430

– 650°C) to form ammonia and isocyanic acid (NHCO) [ 37 ]

. The capital costs of these systems vary with design [ 37 ]

. Urea feedstock is available in solid form or as urea solution in deionized water [ 6 5 ]

. Almost all urea

- to

- ammonia systems use solid urea [ 6 5 ]

. U rea

- to

- ammonia systems typically include dry urea unloading equipment, storage silo, dissolving tank using deionized water, feed tanks, feed pumps, a solution heater

, and a hydrolyzing reactor or decomposition chamber, depending on the type of process used [ 6 5 ]

. Descriptions of two types of these systems are provided.

One of the urea

- derived ammonia production systems converts urea by thermal hydrolysis to ammonia, carbon dioxide, and water vapor

. As shown in Figure 2.11

, t he system consists of urea storage, handling, and dissolvers ; a reactor feed tank ; circulation ; feed pump ; condensate skids ; a hydrolysis reactor ; and ammonia flow control units

. In the urea solution preparation, d ry urea is fed batch

- wise from delivery trucks directly to the dissolver along withChapter 2

– Selective Catalytic Reduction

2

- 39 deionized water

, and the urea solution is stored

. In the solution feed and control system, u rea solution is transferred to the reactor feed tank, and the feed pump meters the urea solution to the reactor, which is heated using steam

. The hydrolysis reactor is a kettle

- reboiler type heat exchanger that operates at 80 psig and at a temperature in the range of 280 to 3 1 0 ° F (140

– 150°C)

. The urea

- to

- ammonia reaction occurs in two steps : the first reaction produces ammonium carbamate (NH 4 COONH 2 ) from urea and water, and the ammonium carbamate breaks down into carbon dioxide and ammonia in the presence of heat

. The temperature of the reactor drives the rate of ammonia production

. Trace amounts of formaldehyde are present when formaldehyde

- conditioned urea is used as feedstock; operating the hydrolysis reactor at a pH of 9 or higher limits formation of urea formaldehyde polymeric resins that can deposit on the reactor [ 6 7 ]

. One of the urea

- derived ammonia production systems decomposes urea to generate ammonia that is fed to the AIG, as shown in Figure 2.12. The system consists of a blower, decomposition chamber, urea storage, chemical pumping system, and process controls

. In the urea storage and pumping system, dry urea from the storage tank is mixed in a solution tank with water and transferred to an aqueous urea solution storage tank

. Filtered ambient air is fed into the decomposition chamber through the use of a blower with automatic dampers to control discharge flow and pressure

. In the chamber, a burner is fired downstream of the dampers, and an aqueous urea solution that is supplied by the storage and pumping system is sprayed into the post combustion gases by inject ors

. The decomposition occurs under a specific temperature and residence time, with the decomposition temperature rang ing from 600 to 1000 ° F (320

– 540°C)

, and the urea is decomposed to ammonia and isocyanic acid

. The outlet ammonia stream from the decomposi tion chamber feeds into the AIG system for the SCR [ 6 8 ].

Figure 2

- 11 : Urea

- Derived Ammonia Production System Using U2A system [ 6 5 ]

Used by permission of R. Salib of URS Washington Division (formerly Washington Group International)

.Chapter 2

– Selective Catalytic Reduction

2

- 40 Figure 2.12: Urea

- Derived Ammonia Production System Using NO x ULTRA System [ 6 8 ]

Used by permission of K.R. Dougherty of Fuel Tech, Inc.
Ammonia Injection
For aqueous

, anhydrous , or urea

- derived ammonia, the ammonia

- air mixture is directed through a flow

- balancing skid to the AIG

, where it is injected under pressure. The flow

balancing skid consists of flow meters and manual valves to ad just the flow to each part of the AIG.

The AIG consists of a network of pipes or lances c onnected in parallel and perfo rated with several holes or nozzles. The lances are placed in a grid formation across the width and height of the ductwork. The lances and holes are sized to distribute the ammonia uniformly into
t he flue gas. The spray angle and velocity of t he injection control the trajec tory of the ammonia

. Injectors are subject to high temperatures and flue gas impingement

, which cause erosion, corrosion, and degradation of structural integrity. Therefore, injec tors are generally constructed of stainless steel and designed to be replaceable. Multiple injec tion zones may be used to increase the distribution of ammonia.
The ammonia can be injected with a low

- or high

- energy system. A low

- energy system uses little o r no pressurized air while a high

- energy system uses large amounts of compressed air or steam to inject and vigorously mix the solution with the flue gas. AIG systems in large boilers typicall y use high

- energy systems. High

- energy systems are more expensiv e to build and operate because they require a larger compressor and a more robust injection system, and consume more electric power.Chapter 2

– Selective Catalytic Reduction

2

- 41 Uniform distribution and mixing with flue gas is critical to maintain desired low levels of ammonia slip. Cold gas flow mod eling and nu merical flow modeling are gener ally performed for the AIG and SCR system to ensure uniform mixing and dispersion before the gases enter the SCR reactor. If duct length is inade quate to ensure thorough mixing or results from the model study indi cate poor gas mixing characteristics, devices such as turning vanes or static gas mixers may be added.
An essential part of an AIG system is the controller used to regulate ammonia injection. Boiler load, inlet NO x , and inlet gas temperatures set t he feed

- forward signal to estab lish the base ammonia injection rate. A feedback signal measuring the SCR outlet NO x concentration is used to trim the base ammonia injection rate.
Although not necessarily required to achiev e high NO x removal efficiencies and low ammonia slip, most SCR installations today employ some type of static mixer to achieve good NH 3 /NO mixing [ 37 ]

. Static mixers provide a more uniform flux of NO x and more uniform temperature and mixing of NH 3 and NO [ 37 ]

. For example, a chiev ing 90 % NO x remo val and a 2

- ppm ammonia slip typically requires NH 3 /NO uniformity l ess than 5 % and perhaps as low as 3 % , as measured on a root mean square basis , and static mixers enable the SCR to achieve these levels [ 37 ]

. T he costs for static mixers vary (e.g., a 500 M W unit may have an installed cost of $7 50,000 and an additional 1 in. w.g. of flue gas pressure drop) [ 37 ]

. One specific type of static mixer is the Delta Wing TM mixer, which consists of an obstruction in the duct, usually a stationary disk or triangular plate, oriented at a slant to the flow direction

. The Delta Wing mixer creates large vortices downstream of the device, which promotes mixing to a more homogenous
gas

. The amm onia injection nozzles are located in the vortex zone immediately downstream of the mixer [ 6 9 ]

. The Delta Wing mixer costs for installing a new SCR typically are less than $ 500,000 and include the capital costs for the mixer and the modeling necessary to d etermine the location for maximum mixing effects. For an existing SCR that was installed without a static mixer, the costs to install the Delta Wing mixer, including capital, modeling, and other retrofit costs, could be up to $1,000,000 [ 70 ]

. Another appro ach that may help retain good NH 3 /NO uniformity is to conduct an annual tuning of the AIG , which can improve the NH 3 /NO mixing (reduces NH 3 /NO “ unmixedness ” by 2

– 5 % on a root mean square basis) [ 37 ]

. Data have show n that the ability of the AIG to achieve good mixing can decline over time and that annual tuning can return the AIG to startup or near

startup mixing uniformity [ 37 ]

. Annual AIG tuning can cost from $30,000 to $50 ,000 depending on the unit size [ 37 ]

. Depen ding on the type of mixer used, annual tuning may not be necessary

, because s ome static mixers combine gas mixing and reagent injection in one application, with no moving parts in the gas stream; this type of mixer avoids much tuning during startup and

com missioning and for annual maintenance [ 7 1 ]

. Use of static mixers and annual tuning can either increase the NO x efficiency at the same ammonia slip level or extend the catalyst life at the same NO x removal efficiency [ 37 ]

. Extending the catalyst life can si gnificantly reduce operating costs , even after accounting for the outsourced tuning costs [ 37 ]

.Chapter 2

– Selective Catalytic Reduction

2

- 42 Catalytic Reduction of NO x The catalytic reduction of NO x in the SCR reactor occurs when the NO x and ammo nia in the flue gas contact the catalyst layers. The catalyst itself is the key component of the SCR system. The catalyst composition, type (honeycomb , corrugated, or plate), and physical properties affect performance, reliability, catalyst quantity required, and cost. However, because the SCR system supplie r and catalyst supplier must guarantee catalyst life and performance, most catalyst characteristics are selected by the SCR system supplier.

2.2.5 SCR System Auxiliary Equipment

SCR Inlet and Outlet Ductw ork
In retrofit installations, new ductwork is req uired to integrate the SCR system with the existing equipment. In high

- dust SCR system s for utility and industrial boilers , the react or is located between the econo mizer outlet and the air heater inlet. In low

- dust SCR systems for utility and industrial bo ilers , the SCR reactor is located between the outlet duct of the particulate control device and the air heater inlet duct. In tail

- end SCR systems for utility and industrial boilers , the ductwork tie

- ins are downstream of the FGD system and also require th e integration of the flue gas reheating equipment.

See Section 2

, Generic Equipment and Devices and Chapter 1

, Hoods, Ductwork, and Stacks

, for more details.
SCR Bypass Duct

Low

- load boiler operations can decrease the temperature at the SCR reactor inlet below the SCR operating range. In addition, startup and shutdown of the boiler causes drastic temperature fluctuations. For these operating conditions, an SCR bypass may (but not necessarily) be required to route the flue gas around the reactor chamber. The bypass prevents catalyst poisoning and fouling during periods when flue gas stream conditions do not meet design specifications for proper SCR operation

. The bypass system also must include zero

leakage dampers to prevent flue gas leakage from poisoning and fouling the catalyst while the SCR is not operating. A bypass system may also be considered for seasonal operation of the SCR system , such as for boilers that would require NO x control during the ozone season ( typically May to September), but not at other times of the year

. An SCR bypass may be needed for cement plant applications. During periods of startup and shutdown, the operating temperatures and constituents in the kiln g ases may affect SCR operating conditions and may cause catalyst plugging or damage.

It should be noted that operational routines can be applied during SCR startup and shutd own that could preclude the need for a bypass , however , particularly for SCRs that o perate year round [56]

. Also, a Haldor Topsoe paper indicates that a bypass is not recommended for reasons including: complicated flue gas duct work, increased risk of dust depositing in horizontal parts, dust precipitation around dampers, and erosion of louver

- type dampers that mayChapter 2

– Selective Catalytic Reduction

2

- 43 result. 17 In fact, many of the SCRs being built in the late 2000’s for compliance with the Clean Air Interstate Rule (CAIR) do not include bypasses. 18 Recently built coal fired power plants in Germany are designed without any S CR bypasses, according to a report by a major engineering services firm. 19

Soot B lower or Sonic Horn

In coal

- fired boilers, soot blower s are usually installed in the SCR reactor to remove particulates that may mask or block active catalyst surfaces and gas passages. Soot blowing helps maintain acceptable flue gas pressure drop in the SCR reactor by keeping the catalyst gas passages free of particulate. Soot blower s also keep the air heater gas passages open and thereby reduce system pressure drop. This is es pecially true for SCR retrofits where the air heater plate spacing is generally narrow, making it more susceptible to fouling or clogging by ammonia

sulfur salts.

Retractable rake

- type soot blower s that use steam or air for blowing are used in SCR designs. The soot blower s are typically located above each catalyst layer. Soot blowing is usually performed on one catalyst layer or part of one catalyst layer at a time. Soot blowing of all the catalyst layers takes 30 minutes to 2 hour s, but is usually done infre quently. In European SCR installations, soot blowing is done approximately once or twice a week [ 7 2 ]. Traveling

- rake steam soot blowers can have installed costs of $120,000

– $160,000 [ 37 ].

Use of sonic horns has risen as an alternative to soot blower s [ 37 ]

. Sonic horns require less preventive maintenance than soot blower s , but they are susceptible to moisture and fly ash, which cause plugging of the horn [ 37 ]

. Sonic horns also cannot damage catalyst through either high

- pressure operation or steam leaks, as can occur with soot blower s [ 37 ]

. The capital cost for sonic horns can be $40,000

– $100,000 for each catalyst layer, depending on the size of the unit [ 37 ]

. Sonic horn operating costs have been reported from $1 /day to < $4/day for each catalyst layer , compared with approximately $41/day for conventional soot blower s [ 37 ]

. Although sonic horns may have some advantages over soot blower s, the demand for high NO x removal efficiency requires extremely clean catalyst and thus the bes t cleaning system (regardless of cost); therefore

, higher costs for soot blower s are justified for many applications [ 37 ]

. For high dust loadings, the cement industry also reports use of soot blowers using heated, dry compressed air and use of sonic horns [94]. Improvements to cement dust cleaning for high dust loading and sticky dust have been made by changing the geometry and operating parameters of the dust blowers [ 112 ].
Large Particle Ash (LPA) Equipment

A significant concern for utility and industri al boiler SCR operation that was not evident in early applications is the role of the accumulation of LPA, also referred to as “popcorn ash,” on

17 Jenson

- Holm, Hans, Lindenhoff, Peter, and Safronov, Sergey. SCR Design Issues in Reduction of NO x Emissions from Thermal Power Plants. Haldor Topsoe A/S, Russia Power, 2007.
18 Rutherford, Scott. Cormetech, Inc. Coal

- Fired Applications in the U.S.

– Challenges and Strategies for Succes sful Operation and Emissions Compliance. VGB Workshop, “Flue Gas Cleaning 2007”. May 22

- 23, 2007, Vienna, Austria. Available on the Internet at http://www.cormetech.com/brochu res/2007_VGB_Conference_Paper.pdf

. 19 Nielsen, Flemming Skovgaard, Danesi, Paolo, and Radhakrishnan, M.V. Modern Boiler Design, BWE. January 2012.Chapter 2

– Selective Catalytic Reduction

2

- 44 catalyst surfaces of high

- dust SCR applications

. LPA is defined as particles that are 4

– 7 mm in characteristic dimension and large enough to lodge in the openings of grid

- or plate

- type catalysts [37 ]

. It is estimated that up to half of SCR units on coal

- fired utility boilers are affected by LPA [8 ]

. LPA is not an issue for natural gas

- fired applications.

The cause or mechanism by which LPA is formed is unknown. A survey of 32 utility boiler operators found 23 had experienced significant plugging of catalyst modules, but there was no clear correlation of design or operating characteristics (e.g., coal type, boiler o r reactor design, SCR cleaning method, or catalyst geometry) with either significant LPA problems or the lack of problems. Without knowledge of how LPA is generated, many facilities mitigate its impacts by removing some of the LPA from the flue gas before it reaches the catalyst

. The most common mitigation method is the use of screens and/or baffles between the economizer exit and the SCR reactor that provide a barrier to the LPA and divert it to an ash hopper . This approach was used by operators of 21 of t he 23 surveyed boilers that experienced LPA problems , and the other 2 were considering adding such equipment. Some of the 9 boilers that were not equipped with screens also had flow or deflector baffles. The frequency of economizer ash hopper evacuation ha s been suggested as a key variable affecting catalyst blockage. Increasing this frequency would be a low

- cost operating change, but the survey of utility boiler operators did not show a correlation between evacuation frequency and catalyst blockage levels [7 3 ]

. The survey report also identified the following recommendations for effective use of screens: (1) orient the screen at an angle to the flue gas flow, or use pleats; (2) maintain at least 50% to 60% open flow area; (3) conduct CFD or physical flow modeling and design the duct and

screen to keep flue gas velocity below 50 actual ft/s (or preferably <45 actual ft/s) and eliminate peaks in velocity; an d ( 4 ) use active cleaning systems

. The open flow area and velocity recommendatio ns are intended to minimize erosion of the screen material by fly ash , which was found to be significant regardless of the screen material used at the surveyed facilities [ 73 ]. The low velocity may also encourage LPA to drop out of the flue gas [37 ]. LPA s creens can be modular to allow replacement, coated or uncoated depending on velocity , and rigid or flexible [37 ]

. Capital costs for a simple rigid screen can be $200,000

– $500,000 for an erosio n

- tolerant design for high flue

- gas velocity and exotic construc tion material [37 ]

. The installed cost of screens for t wo SCRs on t w in boilers that collectively generate 745 MW was $600,000 in 2004 [7 4 ]. A redesign and replacement of this screen along with CFD flow modeling in 2009 cost $806,000 [7 5 , 7 6 ]. Operating cos ts can be a 1

- in. w.g. pressure drop and require an additional $150,000 every 2 years to replace eroded screen panels

. The costs associated with LPA mitigation methods can be recovered by avoiding an outage, by not accelerating an outage, and by maintaining clean catalyst [ 37 ]

. Another option for LPA mitigation involves the use of targeted in

- furnace injection, which models injections strategies to reduce SO 3 formation, and also minimizes slag and fouling. This s lag and fouling control also reduces LPA formation. A suspended slurry of magnesium hydroxide (Mg(OH) 2 ) is used to change the slag formation by traveling into a furnace, becoming superheated and subsequently forming very small particles of magnesium oxide (MgO). The performance of targeted injection, and associated reduction of LPA, has been established in a case study on a 600 MW opposed wall

- fired unit, where successful control of LPA has beenChapter 2

– Selective Catalytic Reduction

2

- 45 demonstrated for more than six years. After succes s fully imple menting the targeted in

- furnace injection

, the facility was able to remove their pre

- existing LPA screens [ 7 7 ]

. Economizer Bypass Duct
Although the SCR reaction occurs withi n a temperature window of 600

– 750°F (3 20

– 400°C), the catalyst for a given applicati on is designe d for a somewhat narrower range: the economizer outlet temperature at normal boiler operating load. Maintaining the flue gas temperature within the required window is essential for optimizing the NO x reduction reaction. When the economizer out let flue gas temperature decreases because the plant is operating at reduced loads, the temperature can be raised using an economizer bypass.
The economizer bypass duct generally has a modulating damper to regulate the amount of hot bypass gas flow to be m ixed with the cooler economizer outlet flue gas. The lower the boiler load, the more this damper opens, thus admitting more hot gas. The economizer outlet duct also needs a modulating damper to provide enough backpressure to allow the required volume of
ga s to flow through the bypass. The main design considerations for an economizer bypass involve maintaining the optimum gas temperature and ensuring uniform mixing of the two gas streams prior to entering the SCR reactor.
Upgraded or New Induced Draft (ID) F an

The new ductwork and the SCR reactor ’ s catalyst la yers decrease the flue gas pres sure. T o maintain the same flow rate through the duct work, additional energy is required. The existing ID fan may be unable to provide the required increase in static pressure. In such cases, an upgraded or new ID fan is installed. This is also true for cement kiln applications due to the high dust loading and high pressure drop across the SCR. The existing fan and motor foundation
may also need modification. Replacemen t involves installation of a new fan or booster fan. In all cases, additional electric power for the ID fan is needed to overcome the additional pressure drop through the SCR system. Based on typical values for the pressure drop through the additional duct work and the catalyst layers, the additional electric power needed ( i.e. , the heat rate penalty) is equivalent to approximately 0.3 % of the plant ’ s electric output for SCR on a utility boiler

. Refer to Equation s 2.57 and 2.58 for estimate s of the total ad ditional electric power needs for the ID fan as well as other equipment in the SCR system

. 2.2.6 Other Considerations

Fuel Source
Utility and i ndustrial boilers use coal, distillate oil, residual oil, natural gas , and a variety of other fuels such as biom ass ( e.g., wood, bark) . The fuel type and grade affects the SCR design

, and therefore, the capital costs of the SCR system. Fuels with high heating value have higher gas flow rates, whic h in turn increase the required SCR reactor size and catalyst volume. Coal

- fueled applications are more costly than oil

- and natural gas

– fired boilers

, due to their higher flue gas flow rates [ 5 3 ]

. The quantity of nitrogen, fly ash

, and pollutants in the flue gas stream varies according to the type and grade of the fuel. This affects the volume of catalyst required

, as well as the catalyst design, composition

, and rate of deactivation. Coal flue gas contains a greater amount of fly a sh,Chapter 2

– Selective Catalytic Reduction

2

- 46 SO 2 , SO 3 , arsenic

, and other trace pollutants than does flue gas from burning oil or natural gas. Natural gas is the cleanest fuel and contai ns the least nitrogen; therefore, burning natural gas results in the least NO x and pollutants in the flue gas s tream. The amount of nitrogen and pollutants in oil

- based fuels sources varies with the grade and type of oil, either refined or residual.

Formation of SO x Sulfur trioxide ( SO 3 ) forms during the combustion of fuels that contain sulfur

, and additional SO 3 i s formed over the SCR catalyst

. S O 3 is an emission concern for low sulfur coal. In cement plants, the fuel sulfur is incorporated into the clinker, however, the raw material sulfur can be a source of SO 2 and SO 3 in the gas stream. SO 3 reacts with ammonia in the flue gas downstream of the reactor (ammonia slip) to form ammonium bisulfate (NH 4 HSO 4 ) and ammonium sulfate [(NH 4 ) 2 SO 4 ] . The amount formed depends on the sulfur content of the fuel

, the amount of ammonia slip , the SCR temperature , and the catalyst formulation and activity

. Ammonium bisulfate condenses as the flue gas stream temperature lowers. It then deposits on the SCR catalyst and downstream equipment such as the air heater, ducts, and fans. Ammonia slip limits are generally im posed as part of the SCR design requirements to avoid impacts on downstream equipment.
There are several methods for limiting the impac t of ammonia

- sulfur salt deposi tion. Soot blowers can be installed between catalyst layers to remove surface deposits by blowing air or steam across the catalyst. Increased acid washing of the air preheater and other equipment may be required to remove deposits. However, more frequent acid washing generates additional
wastewater
, which must be disposed of or treated by the p lant. The sulfur content of the flue gas can be decreased by coal desulfurization processes or fuel switching. Lastly, the flue gas temperature may be raised to prevent condensation.

Elevated SO 3 concentrations raise the acid dew point of the flue gas. Thi s phenomenon potentially leads to more corrosion on the air heater ’ s cold

- end surfaces if the flue gas temperature is below the acid dew point. For Ljungstr ö m © preheaters, the cold

- end baskets of the air heater can be replaced with enamel

- coated baskets to protect against this possibility

. Ammonia sulfates also deposit on the fly ash. Ammonia content in the fly ash greater than 5 ppm can result in off

- gassing

, which would impact the salabil i ty of the ash as a byproduct and the storage and disposal of the ash by landfill [ 51 ]

. Oxidation of SO 2 in the flue gas stream to SO 3 over the catalyst bed create s even more SO 3 (a PM 2.5 precursor compound) and is a n emission concern for SCR control devices only if the sulfur content of coal is sufficiently high

. This can be a concern with SCR applied to boilers that use coal containing greater t han 2% sulfur content or higher

. 20 In some areas of the country, limits for SO 3 of 5 ppm or less have been implemented to control SO 3 plumes emitted from SCRs ( i.e., blue plumes ) [ 37 ]

. Mitigation of SO 3 formation from oxidation of SO 2 on the catalyst has evolved to be of equal importan ce to NO x control in SCR design [ 37 ]

. Catalysts designed for low S O 2 oxidation have been developed, but depending on the catalyst supplier, their u se may increase the catalyst volume needed to meet the sa me NO x removal and ammonia slip

20 U.S. EPA. Regional Haze FIP for New Mexico (San Juan NO x BART), Final Rule. 76 FR 52388 (August 22, 2011).Chapter 2

– Selective Catalytic Reduction

2

- 47 performance and may also alter the catalyst management strategy [ 37 ]

. Another approach to mitigat ing SO 3 formation that has been analyzed is u se of reagents such as sodium, magnesium, or calcium

- based sorbents [ 37 ]

. However, the use of reagents alone to lower SO 3 concentration can rival the costs for annual SCR ammonia or urea reagent support and can exceed catalyst replacement charges , although other sources cite significant plant operation and maintenance cost savings due to eliminating corr osion and fouling of ductwork, ESP s , and the blade housing s of ID fans ; eliminat ing modifica tion or coatings for the air heater; and reducing the heat rate and fuel costs [ 37, 7 8 ]

. Use of low

- SO 2 - oxidation catalyst can be used in combination with reagent

based SO 3 mitigation to lower costs o f reagent

- only approaches to reducing SO 3 [ 37 ]

. Formation of Arsenic Oxide

Arsenic oxides (As 2 O 3 ) formed during combustion of fuel containi ng arsenic cause catalyst deactivation by occupying active pore sites. Formation of arsenic oxides is a concern primarily for SCR units on coal

- fired boilers because coal is the only fuel that may contain arsenic

. Gaseous arsenic oxide can physical ly block active sites by condensing in the pores, and As 2 O 3 can react with oxygen on the vanadium compounds in the catalyst to form As 2 O 5 , which chemically bonds to the site [ 7 9 ]

. Catalyst manuf acturers have developed arsenic

- resistant catalysts that include control on porosit y to minimize pore condensation and addition of compounds such as molybdenum or other materials to mitigate deactivation of active site s [ 7 9 ]

. Deactivation of arsenic

- resistant catalysts occurs at a slower rate

, but the se catalyst s are no t

a rsenic proof [ 7 9 ]

. Arsenic

- resistant catalyst will maintain adequate activity for approximately 14,000 hours of operation versus approximately 5,500 hours for a nonresistant catalyst [ 7 9 ]

. The addition of small amounts of calcium to the fuel has been sh own to be an effective method of controlling arsenic poisoning [ 80 ]

. Limestone (CaC O 3 ) can be injected into the flue gas to generate the solid Ca 3 (AsO 4 ) 2 , which does not deposit on the catalyst and can be removed from the flue gas with a precipitator. Burning coal with higher calcium oxide (CaO) concentration can have the same effect as injecting limestone [ 7 9 ]

. It should also be noted that CaO can be a catalyst poison via formation of C a SO 4 in the catalyst pores, effectively blocking catalyst reactive sites [ 80 ]

. For the cement industry, a significant portion of the dust in cement kiln gases is CaO [94]

. Mercury Oxidation
SCR catalysts have been found to oxidize a significant portion of elemental mercury to oxidized mercury (e.g., elemental mercury ox idiz ed to mercuric chloride [ HgCl 2 ] ). The oxidized
mercury is water soluble and easier to remove in downstream wet FGD units or SO 2 scrubbers. Studies have suggested that the oxidation of elemental mercury by SCR may be affected by catalyst space velocity, reaction temperature, ammonia concentration, catalyst age, and concentration of chlorine in the gas str eam. The type of coal burned and its associated chlorine content is another factor [ 8 1 ]

. Studies on simulated flue gas streams and slip streams from actual electric generat ing units for multiple catalyst types have shown that mercury oxidation is in the ran ge of 80

– 90 % for fresh catalyst and space velocities of approximately 1

, 000 hr

- 1

, but that th e oxidation rate declines to 30

– 8 0 % with increasing space velocity, in the range 4

, 000 hr

- 1 , with other factors such as temperature, ammonia concentration, and so on more prominently influencing the rateChapter 2

– Selective Catalytic Reduction

2

- 48 when the space velocity increases. T hese studies also showed that lower temperatures increase the mercury oxidation rate. The oxidation of mercury was higher at temperatures of approximately 700 ° F (370°C) and were relatively lower at approximately 800 ° F (430°C), which is consist ent with the fact that mercury oxidation to Hg Cl 2 i s greater a t lower temperatures [ 8 1 ]

. In earlier simulated flue gas and slip stream studies, it was noted that the ammonia reagent may suppress or interfere with the oxidation of elemental mercury , especially as the catalyst ages, although ammonia showed little or no effect on oxidation rates in studies on full

scale utility boiler SCR systems [ 8 1 ] . In recent studies, where the NH 3 to NO x ratio was varied between 0 and 0.95, the suppressive nature of ammonia on mercury oxidation has been demonstrated [ 8 2 ]

. In the reac tion process on the catalyst sites, ammonia adsorption takes precedence over the site chlorination, therefore the NO x reduction reaction has precedence over the mercury oxidation [ 4 7 ]

. Because the favored ammonia adsorption on the catalyst sites minimizes the chlorinated sites, the NO x reduction reaction may actually inhibit mercury oxidation [ 4 7 ]

. In several studies, it has been observed that with NO x reduction efficiencies up to 90 % , the mercury oxidation readily occurs [ 4 7 ]

. However, at NO x reduction eff icienci es beyond 90%

, mercury oxidation is greatly reduced [ 4 7 ]

. A higher chlorine concentration enhances mercury oxidation; bench

- scale testing suggests that hydrochloric acid ( HCl ) is an important exhaust gas constituent that provides the chlorine for oxidation of mercury to HgCl 2 across the SCR catalyst [ 8 1 ]

. Coal types

, such as subbituminous coal

, that tend to have lower chlorine levels will have lower HCl concentrations in the ex haust gas than bituminous c oal , which has higher chlorine levels [ 8 1 ]

. Significant
mercury oxidation by SCR catalysts occurs with bituminous coal

, and the oxidation rate is less

cer tain with other types of coals [ 8 1 ]

. In recent studies, HCl gas was injecte d directly into the flue gas and was varied from 0 to 150 ppmv Cl; the data demonstrated increased mercury oxidation across th e SCR as chlorine was increased [ 8 2 ]

. C hlorine appears to have the greatest effect on mercury oxidation in the range of 0

– 50 ppm [ 4 7 ]

. The chlorine content does not appear to affect the NO x r eduction performance of the SCR [ 4 7 ]

. There is a close correlation between mercury oxidation on the catalyst and the SO 2 to SO 3
conve rsion mechanism on the catalyst [ 4 7 ]

. The mercury reaction with HCl and oxygen is considered to be a diffusion

- controlled reaction th at takes place in the gas phase [ 4 7 ]

. The
mercury oxidation rate is faster than the diffusion velocity of mercury through the SCR catalyst due to its molecular weight, whil e the reaction rate of SO 2 to SO 3 conversion is slower than the diffusion veloc ity of SO 2 through the catalyst [ 4 7 ] . As mentioned above in S ection 2.2.1 ( see Catalyst heading), by altering the catalyst to favor mercury oxidation based on this reaction mech anism, the active sites oxidize mercury instead of converting SO 2 to SO 3 [ 4 7 ]

. Higher
mercury oxi dation rates of greater than 90% can be achieved while maintaining low SO 2 to SO 3 conversion rates in the ran ge of less than 0.5% [ 4 7 ]

. The age of the catalyst has also been shown to be a factor in mercury oxidation. One study compared the mercury oxidation rate for new, aged, and regenerated catalyst [ 8 2 ]

. It was determined that there is little to no difference in the mercury oxidation ca pability of th e new and regenerated catalysts [ 8 2 ]

. While the operating conditions during testing of the aged catalyst were different from the new and regenerated catalysts (i.e., slightly higher temperature

, which would decrease mercury oxidation; slightl y higher chlorine content
, which would increaseChapter 2

– Selective Catalytic Reduction

2

- 49 mercury oxidation; and lower space velocity), the data showed lower mercury oxidation for the aged catalyst [ 8 2 ]

. Dioxin/ F uran C ontrol
Dioxins and furans are emitted from combustion processes like waste incineration, burning fuels like wood, coal, or oil, and even from forest fires. Other processes, like chlorine bleaching of pulp and paper, for example, can result in small quantities of dioxins and furans [ 109 ]. The SCR systems for control of dioxins and furans were first used in the late 1990s in Europe and Japan [ 110 ].

Research has shown that c ommonly used metal oxide catalysts used in SCR for reducing NO x emissions (e.g., vanadium and tungsten oxides on a titanium or platinum oxide based substrate) from waste incineration systems can also reduce dioxin and furan emissions

. This means SCR allows for an integrated approach to NO x and dioxin/furan control, since only one pollution control device is needed to r educe the emission levels of both pollutants. The catalytic oxidation has been shown to occur in a temperature range of about 240 ° C to 330 ° C. Dioxin/furan destruction efficiencies of 95 to 98% have been demonstrated [ 109 ].

Retrofit Versus New Design an d Balance of Plant
Retrofit of SCR on an existing unit has higher capital costs than SCR installed on a new system

. The re is a wide range of SCR retrofit costs due to site

- specific factors, scope differences, and site congestion [ 10 ]

. S pecific factors that impact the retrofit costs include the following: [ 10 ]

 T he amount of available space between and around the economizer and air heater;

 C ongestion downstream of the air heater ( i.e., buildings, conveyors, existing ESP s , FGD system, ID fan, or stack) ;

 The age/v intage and manufacturer of the boiler;

 The d esign margin of the existing ID fan (i.e., the need to upgrade or replace fan impellers, replace ID fans, or add booster fans);

 The c apacity, condition, and design margins of the electrical distributio n system;

 The d esign margins of the existing structural steel support systems;

 The p ositive and negative design pressure of the furnace and existing ESP; and

 The n umber, nature, and type of existing items that must be relocated to accommodate the SCR an d associated systems

. While not all of these components are specifically applicable to cement applications, each of these issues have an equal counterp art for cement kiln retrofits.
The primary balance of plant impacts include new ID fans and fan foundati ons or less expensive new booster fans , a new air preheater plus foundations or modifications to the existing air preheater , duct reinforcements, economizer duct bypass or modifications, elevated SCR, ESP reinforcement, and relocating the existing ESP flue gas conditioning system [ 10 ]

. Retrofit costs for cyclone or wet bottom wall

- fired boilers are somewhat higher than retrofit costs for dryChapter 2

– Selective Catalytic Reduction

2

- 50 bottom wall

- or tangentially

- fired boilers [ 4 2 ]. Differential retrofit cost s for an SCR in Germany i s approximately $200 per MMBtu/hr ( $ 20/kW) [ 4 2 ]. However, a large part of the capital costs
are not affected by a retrofit, including those for ammonia storage, vaporization, and injection equipment. The capital costs estimated by the equation s in section 2.4 Cost Analysis are for retrofit of an existing SCR. The cost examples in section 2.5 represent an average retrofit difficulty.
Combustion Unit Design and Configuration
Boiler size is one of the primary factors that determines the SCR system capital cost s. In addition, boiler configuration influences SCR costs. Boiler configurations that split the flue gas flow for two or more air preheaters or particulate removal systems require more than one SCR reactor. Additional reactors substantially increase capita l costs. Boiler operations that have varying operating load, frequent startup s /shutdowns, or seasonal operations require an SCR bypass. Additional ductwork, dampers, and control systems increase the SCR system capital costs. The SCR system may require modi fications to draft fans and/or installation of additional fans. This increases both capital and operating costs of the SCR system. In addition, boiler and duct modifications may be required for implosion protection to accommodate increased draft requiremen ts [ 5 3 ]

. In a cement application

, the kiln type, production rate, dust loading, the presence of catalyst poisons, and the magnitude and variability of raw material sulfur, SO 2 , and SO 3 are all important factors impacting capital and operating costs for SCR systems.
Modeling of the SCR System
Computational fluid dynamics and chemical kinetic modeling are performed as part of the design process for SCR

. In addition to c omputational fluid dynamics and chemi cal kinetic modeling, three

- dimensional physical flow modeling, also referred to as cold flow modeling, is generally required. Cold flow modeling ensures that the flow through the SCR reactor provides adequate residence time, achieves uni form mixing of flue gas and ammonia, minimizes linear velocities to prevent catalyst erosion, and minimizes pressure drop across the catalyst layers. It involves constructing a model of the ammonia injection system, mixing area

, and reactor chamber. In a h igh

- dust configuration, this involves mod eling from the economizer outlet to the inlet of the air heater. Typical model scales range from a 1:10 ratio to a 1:12 ratio for large electric utility boilers.



### 2.3 Design Parameters

SCR system design is a proprietary technology. Extensive details of the theory and correlations that can be used to estimate design parameters such as the required catalyst volume are not published in the technical literature [ 4 2 ]. Furthermore, the design is highly site

- specific. In light of these complexities, SCR system design is generally undertaken by providing all of the plant

- and unit

- specific data to the SCR system supplier, who specifies the required catalyst volume and other design parameters base d on prior experience and computational fluid dynamic s and chemical kinetic modeling [ 33 ]

. For some industrial applications, such as cement kilns where flue gas composition varies with the raw materialsChapter 2

– Selective Catalytic Reduction

2

- 51 used, a slip stream pilot study can be conducted to d etermine whether trace elements and dust characteristics of the flue gas are compatible with the selected catalyst.

This section presents an approach to estimat ing design parameters that are elements in the costing equations used in EPA’s Integrated Pla nning Model (IPM) [ 8 ]

. This section also presents an approach to estimating other design parameters that characterize an SCR system but that are not used directly in the costing procedure [33]. Although this approach is based on SCR data for utility boiler s, it provides sufficient accuracy and detail to be of aid in develop ing capital and annual costs estimates for SCR as applied to industrial boilers and potentially other industrial

sources

. 2.3.1 Boiler Heat Input
The primary cost estimation parameter in the methodology presented in Reference [ 33 ] is the maximum potential heat released by the boiler or heat input rate, Q B , expressed as MMBtu/hr . It is obtained from the high er heating value, HHV, of the fuel in Btu per pound (Btu/lb) and the maximum fuel co nsumption rate in pounds per hour (lb/hr), :

(2.3)

W here :
Q B , = maximum heat rate input to the boiler, MMBtu/hr HHV = higher heating value of the fuel, Btu/lb ṁ fuel = maximum fuel consumption rate of the boiler, lb/hr 1/10 6 = conversion factor of 1 MMBtu/10 6 Btu

Table 2. 5 provides the HHV for various coals that may be used if t he HHV for the specific coal used in the project is not available. 21

Table 2. 5 : High er Heating Values for Various Coals

Type of Coal Energy Content (Btu/lb)
Lignite 5,000

– 7,500

Subbituminous 8,000

– 10,000

Bituminous 11,000

– 15,000

Anthracite 14,000

If the boiler produces electricity, its maximum heat input can be estimated using the boiler net plant heat rate ( NPHR ) , in MMBtu per Megawatt

- hour (MM Btu/ MWh) :

(2.4)

21 Another source of EGU fuels data is http://www.eia.gov/electricity/data/eia923/Chapter 2

– Selective Catalytic Reduction

2

- 52 W here :

B MW = boiler MW rating at full load capacity. NP H R = net plant heat rate, MMBtu/MW h

Note that if NPHR is not known (e.g., a cogeneration unit), a conversion value for coal of 10,000 Btu/kWh (or 10 MMBtu/MW h ) can be used as a reasonable estimate ; a conversion value for petroleum of 11,000 Btu/kWh (11 MMBtu/MWh) and for natural gas of 8,200 Btu/kWh (8.2 MMB tu/MWh) can be used [8 3 ]

. 22 Using this value, the heat input rate, Q B , for a coal

- fired unit is :

(2.5)

Where:

10 = estimated NPHR for coal , MMBtu/MWh.

2.3.2 Heat Rate Factor

The heat rate factor (HRF) is the ratio of actual heat rate of the boiler, in terms of the boiler NPHR in MM Btu/ M Wh, compared to a typical heat rate of 10 MM Btu/ M Wh. The developers of the costing methodology presented in section 2 .4.1 determined that using thi s ratio in the equation for capital costs helped account for observed differences in actual costs for different coal

- fired boilers. To maintain consistency with that approach, the same ratio (i.e., with 10 in the denominator) also has been used in the equa tions for oil and gas fired boilers in section 2.4.1. The NPHR is simply the amount of fuel energy that a boiler consumes to generate 1 MWh of electricity, and is determined based on measurements of the electricity generation and fuel consumption over the same period of time. As noted above, if it is not known for a particular boiler, use 10 MM Btu/ MW h.

( 2 .6)

Where:

HRF = Heat rate factor NPHR = net plant heat rate of the system to be costed, MM Btu/ M Wh 10 = the NPHR that is the basis of the SCR base cost module capital cost

, MM Btu/ M Wh

.

2.3. 3 System Capacity Factor
The total system capacity factor, CF total , is a measure of the average annual use of the boiler in conjunction with the SCR system. CF total is given by:

22 In recent years ( 200 3 to 2011), the average NPHR for coal has increased slightly (likely due to aging of equipment), an d the average NPHR for natural gas has decreased slightly (likely due to the increased use of natural gas fuel and the installation of new equipment).Chapter 2

– Selective Catalytic Reduction

2

- 53 C F total = C F plant CF

SCR (2. 7 )

Where:

CF total = total system capacity factor CF plant = boiler capacity, which is the ratio of the actual quantity of fuel burned annually to the potential maximum quantity of fuel burned annually CF SCR = SCR system capacity factor, which is the ratio of the actual days of SCR operation annually to the total number of plant operating days per year.

For utility boilers, the capacity factor of the boiler, CF plant

, is the ratio of actual electric output annual ly to the potential maximum electric annual output , as shown in Equation 2 .8a:

( 2 .8a)

Where:

B MW = boiler MW rating at full load capacity, MWh. B output = annual actual MW output, MW/year.

Alternatively, f

or industrial and utility boilers, the capacity factor of the boiler, CF plan t
, is the ratio of actual quantity of fuel burned annually to the potential maximum quantity of fuel burned annually in pounds , i.e., in lb/hr × 8,760 hr/yr

. CF plant is given by:

(2. 8 )

Where:

actual ṁ fuel = annual actual fuel consumption rate of the boiler, lb maximum ṁ

fuel = annual maximum fuel consumption rate of the boiler, lb

SCR systems can be operated year

- round or only during the specified ozone season (commonly, May 1 to September 30) . The capacity fac t or for the SCR system, CF SCR

, is the ratio of the actual number of SCR operating days, t SCR , to the total number of plant operating days per
year:

(2. 9 )

Where:

t

SCR

, = actual days of SCR operation annually, daysChapter 2

– Selective Catalytic Reduction

2

- 54 t p lant = actual days of plant (or boiler) operation in a year, days.

2.3. 4 Inlet NO x and Stack NO x
Inlet NO x , represented as NO X in

, is the NO x emission level in the flue gas exit stream from a boiler prior to the SCR system. Note that NO X in also accounts for combustion controls if the boiler is equipped with su c h controls. The inlet NO x emissions level , obtained from analyzing the boiler flue gas stream, is gen erally given in lb/MMBtu of NO 2 [ 33 ].

The stack NO x , represented as NO X out is the required NO x emission limit at the stack outlet. It is generally set by the plant or regulatory limits and also given in lb/MMBtu of NO 2 [ 33 ].

2.3. 5 NO x Removal Efficiency

The NO x removal efficiency, represented as η NOx

, is determined from the inlet or uncontrolled NO x level of the boiler at maximum heat input rate, CF plant = 1.0, and the required stack emission limit. The equation for the NO x removal efficiency is given by:

(2. 10 )

Where:

η NOx = NO x removal efficiency, fraction NO xin = inlet NO x level from the boiler, i.e., inlet NO x rate to the SCR, lb/MMBtu (at

maximum heat input rate, CF plant = 1.0) NO x out = outlet NO x rate from the SCR, lb /MMBtu

The required NO x removal efficiency is one of the most influential parameters on the overall SCR system cost [ 5 3 ]. Typically, the annual average outlet NOx should not be less than 0.04 lb/MMBtu, or at a level that results in a removal efficiency gre ater than 90%, unless a guarantee has been obtained from a vendor. Additionally, if a facility is subject to an outlet limit over a time period shorter than annually (e.g., a 30

- day rolling average), then that value should be used in the calculation of the NOx removal efficiency. If a facility is subject to both an annual limit and a short

- term limit, then the annual limit should be used in the calculation of the removal efficiency. It is noted that 0.05 lb/MMBtu outlet NOx based on a 30 day (boiler operati ng) average should be obtainable by a power plant boiler with an SCR system

. 2.3. 6 NO x Removal Rates

The tons of NO x removed annually (ton/yr) are:

NO x Removed

/ yr = NO xin 

NOx Q B t op / 2

, 000 ( 2

. 11 )

Where:

NO x Removed

/ yr = annual mass of NO x removed by the SCR, tons/yr Q B = maximum heat input rate to the boiler, MMBtu/hrChapter 2

– Selective Catalytic Reduction

2

- 55 t op = operating time per year (CF total x 8760), hr/yr 2000 = conversion factor for lb/ton.

The pounds of NO x removed per hour (lb/hr) are:

NO x Removed

/ hr = NO xin 

NOx Q B ( 2

. 12 )

Where:

NO x Removed

/ hr = hourly mass of NO x removed by the SCR, lb/hr NO xin = inlet NO x of the boiler, lb/MMBtu (at maximum heat input rate,

CF plant = 1.0)
 NOx = NO x removal efficiency of the SCR, expressed as a fraction Q B = maximum heat input rate to the boiler , MMBtu/hr.

2.3. 7 Stoichiometric Ratio Factor

The stoichiometric ratio factor (SRF) indicates the actual amount of reagent needed to achieve the targeted NO x reduction. Typical SRF values are higher than theoretical values due to the complexity of the reactions involving the catalyst and limited mixing. Higher SFR values generally result in increased NO x reduction. The SRF is an important parameter in SCR system design because it establishes the reagent use of the SCR system

. The SRF is defined as:

(2. 13 )

. . . In a design developed by a system supplier, the SRF would be adjusted to account for temperature, residence time, degree of mixing, catalyst activity, and allowable ammonia slip for

a specific boiler. No equation for estimating SRF is available for SCR. The value for SRF in a typical SCR system , using ammonia as reagent, is approximately 1.05 [33]

. This value incorporates design margins for ammonia slip and the small amount of NO 2 in the boiler flue
gas , which requires two moles of NH 3 per mole of NO 2 instead of one mole of NH 3 per mole of NO as shown in Equation 2.1 a

. For an SCR system using urea as the reagent, 0.525 is a typical value for SRF [8].

2.3. 8 Flue Gas Flow Rate

The full

- l oad flue gas flow rate, including the typic al design margin of 5 to 15 % , is used to size the SCR reactors and associated catalyst inventory. This flow rate should be obtained from test data or a combustion calculation.

If flow rate data are not available, an approximation of the flue gas flow rate to each of the SCR reactors

, q fluegas , can be calculated using Equation 2.14

.

(2. 14 )Chapter 2

– Selective Catalytic Reduction

2

- 56 W here :

q fluegas = volumetric flue gas flow rate through the SCR , actual cubic feet per minute

(acfm)
q fuel = base case flue gas volumetric flow rate factor

, ft 3 /min

- MMBtu/hr T = operating gas temperature at the inlet to the SCR , °F

n SCR = number of SCR reactor chambers 700 = temperature at which the base case flow rate factor was determined, °F 460 = conversion from degrees Fahrenheit to Rankine

.

“ B ase case” flue gas flow rate factors per unit of heat input for three types of coals are listed in Table 2.6. These factors were calculated u sing procedures in Reference [8 4 ] for typical coals, typical boiler excess air levels (i.e., 20 %), and typical SCR flue gas conditions (

– 10 in ches w.g

. and 700° F ) . Note that similar flow rates are obtained using the oxygen

- based F

- factors, wet basis in Table 19

- 1 of EPA Method 19 in 40 CFR Part 60, Appendi x A

- 7

.

Table 2.6 : Estimated Flue Gas Volumetric Flow Rate Factors for Various Coals

Coal Type Estimate d value of q fuel (ft 3 /min

- MMBtu/hr)

Eastern Bituminous 484
Powder River Basin 516 Lignite 547

Note that in general, the number of reactors, n SCR , is site specific. One SCR reactor per boiler unit is typically required in small high

- dust system designs. However, two SCR reactors may be needed to treat flue gas from a larger boiler or a boiler equipped with two air preheaters. The system designs devel oped for the base and sensitivity cases of this report use one reactor. Study

- level costs of a two

- reactor system are expected to be similar to the cost of a corresponding one

- reactor system because the catalyst, ammonia, economizer bypass, and ID fan
cost s are essentially identical.

2.3. 9 Space Velocity and Area Velocity
The space velocity

, V space , is defined as the inverse of the residence time and is given by:

(2. 15 )

Where:

V space = the volume tric flow rate divided by the catalyst bed volume

, hr

- 1 ResidenceTime = the time necessary for a volume of flue gas equal to the catalyst bed

volume to pass through the catalyst bed

, hrChapter 2

– Selective Catalytic Reduction

2

- 57 Space velocity is calculated from the experimentally measured flue gas volumetric flow rate at th e reactor inlet, represented as q fluegas , and the reactor

/catalyst volume, represented as Vol catalyst , given by:

(2. 16 )

Where:

Vol catalyst = volume of the reactor or catalyst layers, ft 3

SCR system designers and vendors use the concept of area velocity, V area , to account for the reaction being limited to active catalyst sites. The area velocity is calculated from the specific surface area of the catalyst per catalyst volume, A specific , as follow s :

(2. 17 )

Where:

V area = the space velocity divided by the catalyst pore surface area , ft 3 /ft 2 ·hr A specific = the specific surface area of the catalyst divided by the catalyst volume, ft 2 /ft 3

A specific is given in units of length 2 /length 3

, is sometimes referred to as the contact surface area of the catalyst

, and must be provided by the catalyst manufacturer.

2.3. 10 Theoretical NO x Removal Efficiency
Equation 2. 10 defines the NO x removal efficiency. H owever, in SCR

, NO x removal effi ciency changes with catalyst activation. The following theoretical equation allows for estimation of removal efficiency, η NOx , based on the catalyst activity constant, K catalyst , at a specified time, t [ 33 ]. The theoretical NO x removal efficiency is:

 NOx = SFR (1

- e a ) (2. 18 )

where

(2. 19 )

Where:

K catalyst = constant for catalyst activity , changes over time (t)

Both K

catalyst and A specific are typica lly provided by the catalyst manufacturer.Chapter 2

– Selective Catalytic Reduction

2

- 58 According to this equation, the NO x removal efficiency increases with increasing NH 3 / NO x ratio and decreasing space velocity (i.e., increasing catalyst volume for a given gas flow rate). In addition, the equation shows that as the activity of the catalyst decreases over time, the NO x removal also decreases.

2.3. 11 Catalyst Volume
The theoretical catalyst volume required for the SCR system is based on the factors discussed in Section 2.2

, Process Description. Equation s 2.1 6 , 2.1 8 , and 2.1 9 can be combined

an d rearranged to determine the theoretical catalyst volume [ 56 ]. Substituting the definition of space velocity (Equation 2.1 6 ) into the definition of a ( Equation 2.1 9 )

, and then substituting that into Equation 2.1 8 and solving for the volume of the catalyst gives :

(2. 21 )

An empirical equation was developed in Reference [ 33 ] as a function of several sensitivity variables. The sensitivity variables were determined from catalyst volume estimates obtained from catalyst suppliers for base and sensitivity cases. Adjustment factors for these variables were then developed using regression techniques.
The empirical equation for catalyst volume is given below:

(2. 22 )

where N SCR is the number of SCR reactors and the adjustment factors include:

 NO x efficiency adjustment factor

,  adj :
(2. 23 )

 Ammonia slip adjustment factor, Slip adj , for ammonia slips between 2 and 5 ppm:

(2. 24 )

 NO x adjustment factor for inlet NO x , NOx adj :
(2. 25 )

 Sulfur in coal adjustment factor

, S adj :
(2. 26 )

where S is the sulfur content of the fuel by dry weight fraction.Chapter 2

– Selective Catalytic Reduction

2

- 59  The temperature adjustment factor

, T adj , for gas temperatures other than 700°F (370 ° C) :
(2. 27 )

where T is the temperature of the flue gas at the reactor inlet in degrees Fahrenheit (°F).

2.3. 12 SCR Reactor Dimensions
The cross sectional area of the SCR reactor is sized for the flow rate of the flue gas

, in acfm

, and the superficial velocity. A typical value for the superficial velocity is 16 ft/sec (960 ft/min). Using this value for velocity

, the equa tion for the catalyst cross

- sec tional area

, A catalyst , is given by:

(2. 28 )

Where:

A catalyst = cross

- sectional area of the catalyst, ft 2

The SCR reactor cross

- sectional area

, A SCR , is approximately 15 % greater than the catalyst cross

- sectional area to account to the module geometry and hardware:

A SCR = 1.15 x A catalyst (2. 29 )

Where:

A SCR = cross

- sectional area of the SCR reactor, ft 2

The actual dimensions of the SCR depend on the module arrangement in the catalyst layer. The typical cross

- sectional dimensions of a module are 3.3 feet wide by 6.6 feet long (1 m × 2 m) . Therefore, the SCR plan dimensions are approximately multiples of these dimensions. Depending on the number of modules in width and in length, the SCR reactor may be square or rectangular. For the purposes of this report, the SCR reactor can be treated as a square. The

s creening costs are valid for rectangular SCR reactors as long as the aspect ratio (length divided by width) is not too large. Industry standard aspect ratios are between 1.0 and 1.5. For a square reactor, the length, l , and width, w , are estimated by:

(2. 30 )

Where:

w = width of the SCR reactor, ft

l = length of the SCR reactor, ft

An initia l value for the number of catalyst layers is estimated first. This estimate is then checked by calculating the catalyst height for each la yer. The initial estimate for the number ofChapter 2

– Selective Catalytic Reduction

2

- 60 catalyst layers can be determined from the t otal catalyst volume, the cross

- sectional area of the catalyst, and the estimated height of the catalyst element. A nominal height for the catalyst, h ′ layer , is 3.1 feet [33]

. 23 A first estimate for the number of catalyst layers, n layer , is:

(2. 31 )

Where:

n layer = number of catalyst layers h ′ layer = nominal height of each catalyst layer, ft

This value of n layer is then rounded to the nearest integer. In addition, there must be at least two catalyst layers.
The height of each catalyst layer is calculated using the estimated number of layers. This must result in the height of a catalyst layer, h layer, to be within the standard industry range of 2.5

– 5.0 feet. The height of a catalyst layer is calculated from the following equation:

(2. 32 )

where 1 foot is added to account for space required above and below the catalyst material for module assembly

. The number of catalyst layers calculated above does not include any empty catalyst layers for the future installation of catalyst. An empty catalyst layer is recommended for use with

a CMP . The total number of catalyst layers

, n total , includes all empty catalyst layers that will be installed:

n total = n layer + n empty (2. 33 )

Where:

n total = total number of catalyst layers

n empty = number of empty catalyst layers, included for future catalyst installation

23 The specified nominal value is one value within a range of values for h ′ layer that will give the same values for n layer in Equation 2.31 and h layer in Equation 2.32 for a particular system. The optimum range of values for h ′ layer differs depending on the inlet NO x rate. For example, for a relatively high inlet NO x rate of 0.86 lb/M MBtu, the optimum range of h ′ layer is 3.1 feet to 4.2 feet. The optimum range shifts to smaller values when the inlet NO x rate is lower, but the range is expected to include 3.1 feet for all inlet NO x rates greater than 0.1 lb/MMBtu. Higher values of h ′ lay er than those in the optimum range result in a layer height greater than 5.0 feet, which is outside the standard industry range. Lower values of h ′ layer result in a lower, but still acceptable, value of h layer . Low values of h ′ layer also slightly decrease the amount and cost of catalyst but increase the electricity costs for the ID fan because the number of layers and pressure drop both increase. Thus, a nominal value of 3.1 feet is expected to provide optimum results for most SCR analyses.Chapter 2

– Selective Catalytic Reduction

2

- 61 The height of the SCR reactor, h SCR , including the initial and future catalyst layers, the flow

rectifying layer, space for soot blower s and catalyst loading, but excluding the inlet and outlet ductwork and hoppers

, is determined from the equation:

h SCR = n total ( c 1 + h layer ) + c 2 (2. 34 )

Where:

h SCR = height of the SCR reactor, ft

c 1 = constant based on common industry practice, i.e., 7 ft, ft

c 2 = constant based on common industry practice, i.e., 9 ft, ft

where the constants are based on common industry practice of c 1 = 7 ft and c 2 = 9 ft

. 2.3. 13 Estimating Reagent Consumption and Tank Size
The rate of reagent consumption or mass flow rate of the reagent, reagent , generally

ex pressed as pounds per hour (lb/hr), can be calculated using the inlet NO x in lb/MMBtu and heat input rate

, Q B , in MMBtu/hr.

(2. 35 )

Where :

ṁ reagent = mass flow rate, or consumption rate, of the reagent, lb/hr M reagent = the m olecular weight of the reagent (60.06 pounds per mole [ lb /mole] for urea,

17.0 3 lb /mole for ammonia) M NO x = the molecular weight of NO 2 (46.01 lb /mole).
 NOx = NO x removal efficiency of the SCR, expressed as a fraction

The molecular weight of NO 2 is used because the NO x emissions, NO x in , a re given in lb/MMBtu of NO 2

. For ammonia, the mass flow rate of the aqueous reagent solution,

, is given by:

(2. 36 )

Where:

ṁ sol = mass flow rate of the aqueous reagent solution, lb/hr C sol = the concentration of the aqueous reagent solution, by weight fraction.

The solution volume flow rate, q sol , generally expressed as gallons per hour (gph), is:Chapter 2

– Selective Catalytic Reduction

2

- 62 (2. 37 )

Where:

q sol = solution volume flow rate, gph

= t he density of the aqueous reagent solution, lb/ft 3
7.4805 = conversion factor of 7.4805 gal/1 ft 3

.

The ρ sol is 56.0 lb/ft 3 for a 29% solution ammonia and 71.0 lb/ft 3 for a 50% urea solution at 60°F.
The total volume stored in the tank, or tanks, is based on the volume that the SCR system requires for operating a specified number of days. The volume stored onsite for the number of operating days, t

storage , is:

Vol tank = q sol × t storage × 24 (2. 38 )

Where:

Vol t ank = total volume of aqueous solution stored in the tank(s), gallons (gal)

t storage = number of operating days the SCR is required to operate between solution

delivery, days 24 = conversion factor of 24 hr/1 day.

Note that the tank volume is typically based on full

- load operation, so the capacity factor is not included in Equation 2. 38 . A common on site storage requirement is for 14 days of SCR operation.



### 2.4 Cost Analysis

The cost

- estimating methodology presente d here provides a tool to estimate study

- level costs. Actual selection of the most cost

- effective option should be based on a detailed engineering study and cost quotations from the system suppliers. The costs presented here are expressed in 201 2 dollars ( 201 2 $)

. 24

The cost equations are based on the EPA Clean Air Markets Division ( CAMD ) IPM [ 8 ]

. In the costing method for SCR from the IPM

, the purchased equipment cost

, the direct installation cost , and the indirect installation cost are estimated together. This methodology is different from the EPA Air Pollution Control Cost Manual methodology, which estimates equipment costs and installation costs separately. Due to the limited availability of equipment cost data and installati on co st data, the IPM equa tions for SCR capital costs were not

24 For cost escalation or de

- escalation, one suggested index is the Chemical Engineering Plant Cost Index (CEPCI). More information on CEPCI values and the indexing procedure can be found at http://www.c hemengonline.com/pci

- home . Other cost indexes are also available. For more information on cost escalation or de

- escalation, please refer to the cost methodology chapter in the Cost Manual (Section 1, Chapter 2).Chapter 2

– Selective Catalytic Reduction

2

- 63 reformulated for this analysis

. 25 One d ifference between the IPM methodology and the methodology presented here is that the IPM methodology includes owner’s costs (for owner activities related to en gineering, management, and procurement) and financing mechanisms ( i .e., allowance for funds used during construction [AFUDC] )

. As stated in the cost methodology in this Manual (Section 1, Chapter 2), o wner’s costs and AFUDC costs are capital cost items that are not included in the EPA Control Cost Manual methodology, and thus are not included in the TCI estimates in this section.
Capital cost equations are provided for both coal

- fired and oil

- or gas

- fired units. The capital cost equations are applicabl e to coal

- fired utility boilers and to oil

- or gas

- fired utility boilers at facilities with generating capacity greater than or equal to (≥) 25 MW. Equations are also provided for coal

- fired and oil

- or gas

- fired industrial boilers with a heat input capaci ty greater than approximately 250 MMBtu/hr. The capital costs estimated by the equation represent average retrofit costs; costs for new construction are typically 20

- 30% less than for average retrofits for units of the same size and design . T he TCI equatio ns include a retrofit factor (RF). A retrofit factor of 0. 8 should be used for new construction and a retrofit factor of 1 should be used for average retrofits

. T he equations may overestimate costs for some simple retrofits of existing plants

. For retrofi ts that are more complicated than average, a retrofit factor of greater than 1 can be used to estimate capital cost s provided the reasons for using a higher retrofit factor are appropriate and fully documented

. The SCR system design shown in the discussion below is a high

- dust configuration with one SCR reactor per combustion unit . It uses aqueous ammonia as the reagent with an allowed ammonia slip in the range of 2 to 5 ppm

. 26 T he catalyst is a ceramic honeycomb with an operating life of 3 years at full load operations [ 33 ]

. The cost equations are sufficient for NO Xout emission levels as low as 0.07 lb/MMBtu for bituminous coal and 0.05 lb/MMBtu for both PRB and lignite coal [ 8 ]

. In general, there are differences in capital cost, operating cost, and process risk for high

- dust and tail

- end SCR units [ 5 2 ]. High

- dust units tend to have higher capital costs when there are space constraints while tail

- end units tend to have higher oper ating costs [ 5 2 ]

. Differences in the cost elements related to tail

- end units are pointed out in the discussions for total capital investment and total annual costs in the sections that follow.
The cost information presented in this report is based on using ceramic honeycomb catalyst for the base case. In general, more catalyst volume is required for an SCR system using plate catalyst, although the unit cost of plate catalyst is lower than honeycomb. Thus, any difference in cost is expected to be within the accuracy of a study

- level cost estimate.
The annual cost procedures in the IPM categorize annual cost element s as either “fixed O&M” costs or “variable O&M” costs. These elements are reclassified in this report as variable

25 The EPA CAMD IPM methodology for estimati ng capital costs is based on an engineering and design firm’s in

house databases of actual SCR projects. The documentation indicates that the current industry trend is to retrofit high

- dust hot

- side SCR, and cold

- side tail

- end SCRs encompass a small minori ty of units and were not considered in the evaluation. Thus, the SCR cost equations are likely most representative of high

- dust SCR, and qualitative differences in equipment and costs are noted in the text for tail

- end units.
26 While the EPA CAMD IPM cost method is based on use of urea

- derived ammonia injection, the system design and example problems here are based on use of aqueous ammonia because the majority of SCR use ammonia as the reagent. This approach may slightly overestimate the capital costs for an ammonia

- based system, as discussed in the “Reagent Production, Storage, and Vaporization” subsection of section 2.2.4.Chapter 2

– Selective Catalytic Reduction

2

- 64 direct annual costs

, semivariable direct annual costs , or indirect annual costs to be consistent with EPA Control Cost Methodology. The proced ures in this report also include capital recovery, which is not included in the IPM procedures.

2.4.1 Total Capital Investment

Total capital investment (TCI) includes direct and indirect costs associated with purchasing and installing SCR equipment. Cost s include the equipment cost for the SCR system itself, the cost of auxiliary equipment, direct and indirect installation costs, additional costs due to installation such as asbestos removal, costs for buildings and site preparation, offsite facilities, la nd , and working capital . In general, SCR does not require buildings, site preparation, offsite facilities, land, and working capital. A more detailed discussion of capital costs can be found in Section 1, Chapter 2 of this Manual. The total project cost or TCI for the SCR is based on the approach used by EPA CAMD in the Integrated Planning Model [ 8 ] , and this approach includes both the direct capital costs and the indirect capital costs

. The methods presented in sections 2.4.1.1 and 2.4.1.2 for utility boil ers are identical to the methods in v5.13 of the IPM , except that two elements have been excluded , as noted above . The IPM does not include methods for estimating impacts to industrial boilers. Thus, the methods presented in sections 2.4.1.3 and 2.4.1.4 fo r industrial boilers are based on modified IPM equations; the equations were modified by replacing electricity production ratings with the corresponding typical boiler heat input capacities, as calculated using typical NPHRs, and assuming that SCR costs fo r industrial boilers and utility boilers that have the same heat input capacity would be the same. The capital cost equations included in the manual reflect a process contingency of 5

- 10 percent and a project contingency of 15 percent.

The SCR costs and the balance of plant costs are impacted by the unit’s elevation with respect to sea level. These cost calculations have been developed for SCR systems located within 500 feet of sea level. For SCR systems located at higher elevations, the base SCR uni t cost and balance of plant cost should be increased based on the ratio of the atmospheric pressure between sea level and the location of the system, i.e., atmospheric pressure at sea level divided by atmospheric pressure at the elevation of the unit [8]

. The elevation factor is calculated as follows:

(2. 39 a )

W here :

ELEV F = elevation factor P 0 = atmospheric pressure at sea level, 14.7 pounds per square inch absolute ( psia ) P ELEV = atmospheric pressure at elevation of the unit , psia ( see Table 2.7 for

atmospheric pressures for various elevations)

.

Table 2.7 presents atmospheric pressures for elevations up to 6000 feet above sea level.Chapter 2

– Selective Catalytic Reduction

2

- 65 Table 2.7. Atmospheric Pressure at Different Elevations.
Elevation above sea level, ft Atmospheric pressure, psia 0 14.7

500 14.4

1000 14.2

1500 13.9

2000 13.7

2500 13.4

3000 13.2

3500 12.9

4000 12.7

4500 12.5

5000 12.2

6000 11.8

Alternatively, P ELEV can be calculated using equation 2.39b [ 108 ]:

(2.39b)

Where:

P ELEV = atmospheric pressure at elevation of the unit, psia

h = altitude, feet

2.4.1.1 Utility Boilers (Coal

- fired)
Utility, c oal

- fired units

. The capital cost equation for coal

- fired units (applicable for ≥25 MW) is:

(2. 40 )

W here :

TCI = total capital investment for a SCR on a coal

- fired boiler, $ SCR Cost = cost of the SCR, $Chapter 2

– Selective Catalytic Reduction

2

- 6 6 RPC = reagent preparation cost, $ APHC = air pre

- heater cost, $ BPC = balance of plant costs, $

This TCI calculation includes a factor of 1.3 to estimate engineering and construction management costs, labor adjustment for installation (e.g., per diem and premium for work shifts of 10 hr)

, and contractor profit and fees. ( For retrofits that are more complicat ed than average, the terms SCR, APHC, and BPC would be adjusted with a retrofit factor of greater than one.)

SCR costs, utility, coal

- fired units ≥ 25 MW

. The capital cost s for the SCR base unit includes costs for the inlet ductwork, the reactor, and the bypass equipment [ 8 ]

. The SCR costs are calculated as follows:

( 2

. 41 )

W here :

SCR Cost = S CR unit costs, $ 2 70 ,00 0 = constant in the equation NRF = NO x removal factor (η NOx /80) B MW = electric generating capacity of the unit supplied by the boiler, MW HRF = heat rate factor CoalF = coal factor ( C oal F=1 if bituminous; C oal F=1.05 if PRB; C oal F=1.07 if Lignite) ELEV F = elevation factor

. RF = retrofit factor (RF = 0.8 for new co nstruction; RF = 1 for retrofits of average difficulty )

The NO x Removal Factor, NRF, is an adjustment factor that helps the equations more accurately reflect the actual costs in the database ; it is expressed as the NO x removal efficiency (η NOx ) divided by 80

. The CoalF is 1 for bituminous coal, is 1 .05 for P owder R iver B asin ( PRB ) coal, and is 1.07 for lignite coal.
Reagent Preparation costs, utility, coal

- fired units ≥25 MW . The costs for equipment to prepare reagents for injection into the SCR are based on the NO x removal rate. As noted in previous sections, ammonia (either aqueous or anhydrous) is typically used as the reagent for SCR

. The RPC equation is applicable for all types of reagent systems because it reflects the actual mix of types of reagent s ystems in the underlying database. As a result it likely slightly overstates costs for anhydrous ammonia systems, and slightly understates costs for urea to ammonia systems

. The reagent preparation costs are calculated as follows:

( 2

. 42 )
W here :

RPC = Reagent preparation cost, $ 4 90,000 = constant in the equation NO xin = inlet NO x level from the boiler, i.e., inlet NO x rate to the SCR,

lb/MMBtuChapter 2

– Selective Catalytic Reduction

2

- 67 NPRH = net plant heat rate, MMBtu/MWh
η NOx = NO x removal efficiency, fraction RF = retrofit factor (RF = 0.8 for new construction; RF = 1 for retrofits of average difficulty )

Air Pre

- Heater Modification costs, utility, coal

- fired units ≥ 25 MW . Air pre

- heater modification costs are included only where SO 3 control is necessary. An air pre

- heater modification is necessary for the control of SO 3 for boilers that burn bituminous coal where the SO 2 content in the outlet stream from the boiler is 3 lb/MMBtu or greater. Such modifications can include the use of steels resistant to corrosion

, sootblowers and nozzles specifically designed to minimize SO 3 formation. If lower sulfur content coal types are used, then no air pre

- heater modification is needed. The air pre

- heater modification costs are calculate d as follows:

( 2

. 43 )

W here :

APHC = Air pre

- heater cost, $ 69,000 = constant in the equation AHF = air heater factor (AHF =1 if bituminous coal and SO 2 ≥3 lb/MMBtu ; if not true,

AHF = 0 )

. RF = retrofit factor (RF = 0.8 for new construction; RF = 1 for retrofits of average

difficulty)

The AHF is 1 for bituminous coal and where the SO 2 content of the coal is 3 lb/MMBtu or greater

. If the boiler burns other coal types, then the AHF is 0 and the air pre

- heater term drops out of the overall TCI equation for the SCR system

. Balance of plant costs, utility, coal

- fired units ≥ 25 MW . The BPC include cost items such as ID and booster fans, piping, and auxiliary power modifications necessary for the SCR unit [ 8 ]

. The BPC are calculated as follows:

(2

. 44 )

W here :

BPC = Balance of plant cost , $ 4 60,000 = constant in the equation ELEV F = elevation factor. RF = retrofit factor (RF = 0.8 for new construction; RF = 1 for retrofits of average

difficulty )

2.4.1.2 Utility Boilers (Oil

- and Gas

- fired)Chapter 2

– Selective Catalytic Reduction

2

- 68 Utility, o il

- and g as

- fired units ( ≥25 MW to 500 MW)

. The capital cost equation for oil

and gas

- fired units

, applicable to ≥25 MW to 500 MW

, is:

(2. 45 )

W here :

TCI = total capital investment for a n SCR unit on an oil

- fired or gas

- fired boiler, $ 8 0 ,000 = installed cost of an SCR system in 20 1 2 $ for a 200 MW oil

- or gas

- fired boiler

,

$/ M W B MW = electric generating capacity of the unit supplied by the boiler, MW ELEV F = elevation factor. RF = retrofit factor (RF = 0.8 for new construction; RF = 1 for retrofits of average difficulty )

The ELEVF was applied to the TCI equation for oil

- and gas

- fired units, however, because the ELEVF would apply only to the base SCR cost and the BPC but not to the RP costs, or to the APHC if applicable, including the factor may somewhat overestimate the TCI

. For utility oil

and gas

- fired units, unlike t he utility coal

- fired units, the TCI equation is a single equation for all capital costs that does not allow discrimination in applying the factor.
Utility, o il

- and g as

- fired units (>500 MW)

. For oil

- and gas

- fired units >500 MW, the normalized costs in $ /k

W are assumed to be equivalent to the costs for a 500 MW boiler unit. Thus , the equation for units > 500 MW reduces to:

(2. 46 )

With respect to high

- dust and tail

- end SCR units, h igh

- dust units typically require larger catalyst volu me that increase s capital costs. Tail

- end units require less catalyst volume and therefore lower capital costs, due to minimal ash and catalyst poisons in the flue gas following the ESP and wet scrubber

. A rule of thumb for SCR catalyst volume is high

- dust units (on cyclone

- fired boilers) require approximately 1.5 m 3 /MW

, and tail

- end units require less than half the catalyst volume of a high dust unit (or less than 0.75 m 3 /MW) [8 5 ]

. The lower catalyst volume for tail

end units helps reduce init ial catalyst capital cost (and catalyst replacement operating costs) [ 62 ]. The capital costs for tail

- end SCR units must include the equipment cost for reheating the flue gas. Reheating may be conducted using steam coils or natural gas firing. Capital cost s for these reheating options are similar, however steam supply piping, supports, and valves may increase the steam coil reheating capital costs [ 62 ]. In a case study for a tail

- end SCR on a 600 MW burning bituminous coal, one source cites SCR capital cost s of $205 million for an SCR with steam coil reheating and $205 million for an SCR with a natural gas burner (2008$) [ 62 ].
As noted earlier, applying the ELEVF to the TCI equation for utility oil

- and gas

- fired units may overestimate the costs.Chapter 2

– Selective Catalytic Reduction

2

- 69 2.4.1. 3 Ind ustrial Boilers (Coal

- fired)

Industrial, c oal

- fired units . The capital cost equation for coal

- fired units (applicable for ≥250 MMBtu/hr) uses the utility equations to estimate the industrial boiler SCR costs. Use of the utility equations may overestimate t he costs for industrial boilers since current retrofits of utility boilers are likely more complex than for industrial boilers . The capital cost equation is:

(2. 47 )

W here :

TCI = total capital investment for a SCR on a coal

- fired boiler, $ SCR Cost = cost of the SCR, $ RPC = reagent preparation cost, $ APHC = air pre

- heater cost, $ BPC = balance of plant costs, $.

This TCI calculation includes a factor of 1.3 to estimate engineering and construction management costs, labor adjustment for installation (e.g., per diem and premium for work shifts of 10 hr) , and contractor profit and fees.

SCR costs, industrial, coal

- fired units ≥250 MMBtu/hr

. The capital cost s for the SCR b ase unit includes costs for the inlet ductwork, the reactor, and the bypass equipment [8]. The SCR costs are calculated as follows:

( 2

. 48 )

W here :

SCR Cost = SCR unit costs, $ 2 70,000 = constant in the equation
NRF = NO x removal factor ( 

### η NOx /80) Q B = maximum heat rate input to the boiler, MMBtu/hr CoalF = coal factor ( C oal F=1 if bituminous; C oal F=1.05 if PRB; C oal F=1.07 if Lignite) ELEV F = elevation factor

. RF = retrofit factor (RF = 0.8 for new construction; RF = 1 f or retrofits of average difficulty )

The NO x Removal Factor, NRF, is an adjustment factor that helps the equations more accurately reflect the actual costs in the database ; it is expressed as the NO x removal efficiency (η NOx ) divided by 80 . The CoalF is 1 for bituminous coal, is 1 .05 for P owder R iver B asin ( PRB ) coal, and is 1.07 for lignite coal.
Reagent Preparation costs, industrial, coal

- fired units ≥250 MMBtu/hr . The costs for equipment to prepare reagents for injection i nto the SCR are based on the NO x removal rate. As noted in previous sections, ammonia (either aqueous or anhydrous) is typically used as the reagent for SCR. The RPC equation is applicable for all types of reagent systems because itChapter 2

– Selective Catalytic Reduction

2

- 70 reflects the actual mix of types of reagent systems in the underlying database. As a result it likely slightly overstates costs for anhydrous ammonia systems, and slightly understates costs for urea to ammonia systems. The reagent preparation costs are calculated as follows:

( 2

. 49 )

W here :

RPC = Reagent preparation cost, $ 4 90,000 = constant in the equation NO xin = inlet NO x level from the boiler, i.e., inlet NOx rate to the SCR, lb/MMBtu
η NOx = NO x removal efficiency, fraction RF = retrofit factor (RF = 0.8 for new construction; RF = 1 for retrofits of average difficulty )

Air Pre

- Heater Modification costs, industrial, coal

- fired units ≥250 MMBtu/hr . Air pre

heater modification costs are included only where SO 3 control is necessary. An air pre

- heater modification is necessary for the control of SO 3 for boilers that burn bituminous coal where the SO 2 content in the outlet stream from the boiler 3 lb/MMBtu or greater. If other coal types are used, then no air pre

- heater modification is needed. The air pre

- heater modification costs are calculated as follows:

( 2

. 50 )

W here :

APHC = Air pre

- heater cost, $ 69,000 = constant in the equation CoalF = coal factor ( C oal F=1 if bituminous; C oal F =1.05 if PRB; C oal F=1.07 if Lignite) AHF = air heater factor (AHF =1 if bituminous coal and SO 2 ≥3 lb/MMBtu ; if not true,

AHF = 0 ) RF = retrofit factor (RF = 0.8 for new construction; RF = 1 for retrofits of average difficulty )

The AHF is 1 for bituminous coal and where the SO 2 content of the coal is 3 lb/MMBtu or greater. If the boiler burns other coal types, then the AHF is 0 and this term drops out of the overall TCI equation for the SCR system

. Balance of plant costs, industrial, coal

- fired units ≥250 MMBtu/hr . The BPC include cost items such as ID and booster fans, piping, and auxiliary power modifications necessary for the SCR unit [8]. The BPC are calculated as follows:

(2

. 51 )Chapter 2

– Selective Catalytic Reduction

2

- 71 W here :

BPC = Balance of plant cost
, $ 4 60,000 = constant in the equation ELEV F = elevation factor. RF = retrofit factor (RF= 0.8 for new construction; RF= 1 f or retrofits of average difficulty )

2.4.1. 4 Industrial Boilers (Oil

- and Gas

- fired)

Industrial, oil

- fired units (≥2 75 to ≤ 5, 50 0 MMBtu/hr) . The capital cost equation for oil

fired industrial boilers was developed by modifying equation 2.4 5 . The 200 MW rating of the generating unit for the base utility boiler in equation 2.4 5 was converted to a heat input capacity of 2,200 MM Btu/ hr for an oil

- fired boiler by using equation 2.4 with a NPHR of 11,000 Btu/kwh. Similarly, dividing the base cost of $80/kW in equation 2.4 5 by the NPHR of 11,000 Btu/kwh gives an estimated cost of 7,270 $/MMBtu/hr. The resulting equation, applicable to oi l

fired industrial boilers with a heat input capacity of ≥275 to ≤5, 5 00 MMBtu/hr, is:

(2. 52 )

W here :

TCI = total capital investment for SCR on an oil

- fired boiler, $ 7,270 = estimated installed cost of an SCR unit in 201 2 $ for a n oil

- fired boiler that has a

2,200 MMBtu/hr rating at full load capacity, $/MMBtu/hr Q B = maximum heat rate input to the boiler, MMBtu/hr ELEV F = elevation factor RF = retrofit factor (RF = 0.8 for new construction; RF = 1 for retrofits of average

difficulty )

As noted earlier, applying the ELEVF to the TCI equation for industrial oil

- fired units may overestimate the costs.

Industrial, g as

- fired units (≥2 05 to ≤ 4,10 0 MMBtu/hr ) . The capital cost equation for gas

fired industrial boilers wa s d e veloped by modifying equation 2.44 in a manner similar to that described above for oil

- fired boilers, except that the NPHR used in the conversions was 8,200 Btu/kwh

. The resulting equation , applicable to gas

- fired industrial boilers with a heat input capacity of ≥2 05 to ≤ 4,1 00 MMBtu/hr , is:Chapter 2

– Selective Catalytic Reduction

2

- 72 ( 2. 53 )

W here :

TCI = total capital investment for SCR on a gas

- fired boiler, $ 9,760 = estimated installed cost of an SCR unit in 201 2 $ f or a gas

- fired boiler that has a 1,640 MMBtu/hr rating at full load capacity, $/MMBtu/hr Q B = maximum heat rate input to the boiler, MMBtu/hr ELEV F = elevation factor RF = retrofit factor (RF = 0.8 for new construction; RF = 1 f or retrofits of average difficulty )

As noted earlier, applying the ELEVF to the TCI equation for industrial gas

- fired units may overestimate the costs.

Industrial, o il

- fired units (>5, 5 00 MMBtu/hr ) . For oil

- fired industrial boilers >5, 5 00 MMBtu/hr, the normalized costs in $/MMBtu/hr are assumed to be equivalent to the costs for a 5, 5 00 MMBtu/hr boiler unit. Thus, the equation for units >5, 5 00 MMBtu/h r reduces to:

(2. 54 )

As noted earlier, applying the ELEVF to the TCI equation for industrial oil

- fired units may overestimate the costs.

Industrial, g as

- fired units (> 4,1 00 MMBtu/hr) . For gas

- fired industrial boilers > 4,1 00 MMBtu/hr, the normalized costs in $/MMBtu/hr are assumed to be equivalent to the costs for a 4,1 00 MMBtu/hr boiler unit. Thus, the equation for units > 4,1 00 MMBtu/hr reduces to:

(2. 55 )

As noted earlier, applying the ELEVF to the TCI equation for industrial gas

- fired units may overestimate the costs.

2.4. 2 Total Annual Costs

Total annual costs (TAC) consist of direct costs, indirect costs, and recovery credits. Direct annual costs are those proportional to the quantity of waste gas processed b y the control system. Indirect ( fixed) annual costs are independent of the operation of the control system and would be incurred even if it were shut down. No byproduct recovery credits are included because there are no salvageable byproducts generated from the SCR [ 8 6 ]. Each of these costs is discussed in the sections below. A more detailed dis cussion of annual costs can be found in Section 1, Chapter 2 of this Cost Manual.Chapter 2

– Selective Catalytic Reduction

2

- 73 Design parameters are estimated using the maximum annual heat input rate of the boiler to ensure adequate sizing of the SCR system. Annual costs are calculated using the annu al heat input rate of the boiler and SCR system using CF total

. . This ensures that annual costs are based on the actual operating conditions rather than the design case. No escalation of annual costs is included in this procedure in order to be consistent with the cost methodology followed in this Cost Manual as described in Section 1, Chapter 2.

Direct Annual Costs

Direct annual costs

, DAC

, include variable and semivariable costs. Variable direct annual costs account for purchase of reagent and electrical power. Semivariable direct annual costs include operating and supervisory labor cost, maintenance cost, and catalyst replacement cost. These costs are discussed individuall y below. Equations for these variable cost items were derived in Reference [ 8 ].
Operating costs also result from small decreases in boiler efficiency due to opera tion of the economizer bypass. The economizer bypass operation depends on the flow rate of gas bypassed at full and partial loads and the boiler ’ s capacity factor. Another operat ing cost is incurred for the steam or electric power used for compressed air, as required for the relatively infrequent operation of soot blower s. These operating costs are generally small and site

- specific. Therefore, they are not discussed in this report.

(2. 56 )

Operating and Supervisory Labor
In general, operation of an SCR system require s only minimal, operating or supervisory labor. The SCR react or is a stationary device with no moving parts. Further, the SCR system incorporates only a few pieces of rotating equ ipment (e.g., pumps, motors ). The IPM [8] estimates operating labor time as 4 hours per day.

Maintenance

The annual maintenance labor an d material cost in dollars per year ($/yr), including nozzle tip replacement for the injecto rs, is assumed to be 0.5 % of the TCI in dollars [ 7 a , 8 ] 27 :

Annual Maintenance Cost = 0.005 x TCI (2. 57 )

27 Reference [8] applies the 0.5% factor for units smaller than 300 MW and applies 0.3% for larger units, and the factor is applied to the “Base Module” cost rather than the TCI (i.e., the equipment and installation cost before adding 30% for engineering and construction management, labor adjustment, and contractor profit and fees). Since the capital cost estimating procedure for oil

- and gas

- fired units does not include estimation of a Base Module cost, the procedures in this report use the more conventional approach of scaling the annual maintenance cost from the TCI.Chapter 2

– Selective Catalytic Reduction

2

- 74 Reagent Consumption

The annual cost of reagent purchase in $/yr is estimated using the reagent volume flow
rate
, q sol , the operating time per year, t op , and the cost of reagent in dollars per gallon, C ost reag : Annual Reagent Cost = q sol x Cost reag x t op (2. 58 )

where q sol is in gallons per hour (gph ). The example here is for use of an aqueous ammonia solution , i.e., q sol , h owever, if the cost of the ammonia is available in $/lb of ammonia, then the equation to estimate reagent costs would be based on m reagent from Eq uation 2. 3 5 instead of q sol

. The operating time per year, t op , is estimated using the capacity factor, CF total :

(2. 59 )

Utilities
Power consumption for utility boilers

. The electrical power consumption

, P

, in k W is estimated for SCR equipment, ammonia vaporization, water vaporization, and additional ID fan power [ 33 ]. The total additional auxiliary power required is estimated using Equation 2. 60 [8] :

(2. 60 )

Where:

P = electrical power consumption of the SCR system , kW 1,000 = conversion factor for 1,000 kW/MW 0. 00 56 = adjustment or scaling factor, dimensionless

For oil

- and gas

- fired boilers, replace the coal factor with “1”.
Power consumption for industrial boilers

. The electrical power consumption, P , in k W is estimated fo r SCR equipment, ammonia vaporization, water vaporization, and additional ID fan power [ 33 ]. It is estimated by converting Equation 2. 60 to use the heat input to the boiler instead of the electric output as shown in Equation 2. 61 :

(2. 61 )

The annual cost of electricity is estimated from the equation:

Annual Electricity Cost = P x Cost elect x t op (2. 62 )

Where:

Cost elect = cost of electricity

, dollars per kWh ($/kWh).

In general, the power consumption for operating a high

- dust SCR is lower than for low

- dust SCR (data are not available for a tail

- end unit). A relative comparison of power consumption for high

dust and low

- dust SCR is shown in Table 2. 8

. As an example, one source cites a total powerChapter 2

– Selective Catalytic Reduction

2

- 75 consumption of 3,500 kW for a high

- dust SCR and 7,000 kW for a low

- dust SCR on a 440 MW coal

- fired boiler, showing that the power consumption for a low

- dust unit is twice that of the high

- dust unit [ 8 5 ].
Depending on the site configuration and space constraints, application of high

- dust SC R to a boiler where there is significantly large retrofit costs may actually favor use of a tail

- end unit with lower capital costs and more simple retrofit even though operating costs are higher [ 8 5 ]

.

Table 2. 8 : Comparison of Power Consumption for High

- Dust and Low

- Dust SCR [ 8 5 ]

Power component High

- dust SCR Low

- dust SCR Induced draft fans, kW (accounts for largest portion of power needs, at 80 to 90% of total power) Lower Higher (approximately 2x higher)

Ammonia system power, kW Higher (approximately 20% higher) Lower Dilution air blower, kW (second largest portion) Higher (approximately 20% higher) Lower

Dilution air heaters, kW Higher (approximately 25% higher) Lower Ammonia pump, kW Higher Lower Seal Air fans, kW Same Same Electrical and control power consumption, kW Lower Higher (approximately 2x higher) Total power consumption, kW Lower Higher

Catalyst Replacement
The catalyst life is a function of the catalyst activity and ammonia slip. As the catalyst activity decreases with time, the ammonia slip increase s until it reaches the design limit and new catalyst must be added. Catalyst life is usually specified when purchasing the catalyst. For the most common SCR design, the high

- dust SCR, a catalyst layer is typica lly guaranteed for 16,000

– 24,000 operating hours based on information from catalyst vendors. For oil

- and gas

- fired units, the SCR catalyst life is assumed to be 40

, 000 hours , and the catalyst life for some gas

- fired units has been reported to be up to 60 ,000 hours

.
Two methodologies for estimating the annual catalyst replacement cost are presented in this chapter. One methodology is based on estimating the total volume of catalyst, the total number of catalyst layers, the number of layers replaced annual ly, and the future worth of the catalyst. This cost methodology assumes a gua ranteed catalyst life of 24,000 hours or approximately 3 years for close to full time operation

. The second methodology is an empirical equation that is part of the S&L cost metho dology employed for power plants in the IPM [8].

Under catalyst replacement cost methodology 1 , i f the SCR does not have an empty catalyst layer, one approach is to replace all of the catalyst layers at the end of 24,000 operating hours. This very conservative assumption (i.e., likely overestimates the control cost ) has beenChapter 2

– Selective Catalytic Reduction

2

- 76 used in the SCR costs developed in the References [ 86 ] and [ 87 ]. If the SCR includes a spare catalyst layer, then only one catalyst layer is replaced at the end of 24,000 hours) . Most SCR designs include a spare catalyst layer. The cost for catalyst replacement in all the SCR reactors for a given boiler is given by:

Catalyst Replacement Cost = n SCR Vol catalyst CC replace

R layer (2.63 )

Where:

Catalyst Replacement

Cost = cost to replace the SCR catalyst, either replacing all catalyst or

replacing 1 layer at a time, $ Vol catalyst = volume of the r e actor or catalyst layers, ft 3

n SCR = number of SCR reactor chambers CC replace = cost of catalyst

, dollars per cubic foot ($/ft 3 ) R layer = catalyst replacement factor ( R layer = 1 for full replacement and R layer =

n layer for replacing 1 layer at a time)

The catalyst cost, CC replace , should reflect the current costs for the c atalyst, including the costs associated with installing the new catalyst and removing and disposing of the old catalyst

. Where a plant intends to use only regenerated catalyst, the catalyst cost used in equation 2.63 should reflect the current cost of regenerated catalyst, which is typically lower than that for new catalyst. For situations where a plant may use regenerated catalyst and periodically purchase new catalyst, then a weighted average of current costs for new and regenerated catalyst should b e used in equation 2.63.
Because the catalyst is replaced every few years, the annual catalyst cost for all reactors is a function of the future worth of the catalyst, FWF , and is given by:

(2. 64 )

where FWF is the future worth facto r. Future worth is used because the annual catalyst replacement cost is accrued starting in the first year of operation, while catalyst replace ment purchases occur every few years. To account for the time valu e of money, the FWF amortizes the catalyst cos t over the years preceding the actual catalyst purchase [ 88 ]. Because the money is allocated in advance of the purchase, the sum of the annual catalyst replacement costs is less than the purchase price of the catalyst. The future worth factor, FWF is given by:

(2. 65 )

Where:

i = interest rate, fractionChapter 2

– Selective Catalytic Reduction

2

- 77 Y = term, years.

The term, Y

, is given by the equation:

(2. 66 )

Where:

h catalyst = operating life of the catalyst

, hours h year = number of hours per year the SCR is operated , hr/yr

The value of Y estimated from the equation is then rounded to the nearest integer.

Under catalyst replacement cost methodology 2 , t he cost for catalyst replacement and disposal for a given boiler is part of the S&L cost methodology employed for power plants in this chapter given by [8] :

(2. 6 7 )

Where:

Annual Catalyst Replacement Cost = cost to replace the SCR catalyst, $ /yr CC replace = cost of catalyst

, dollars per cubic meter ($/ ft 3 ) 35. 3 = conversion factor for $/ft 3 to $/m 3

.

Because high

- dust units typically require larger catalyst volume, the replacement costs for the catalyst are also higher. Tail

- end units require not only less catalyst volume but also less frequent catalyst replacement, due to minimal ash and catalyst pois ons in the flue gas at this point in the equipment train. Lower levels of fly ash and catalyst poisons in the flue gas increase the catalyst life and decrease operating costs related to replacement [ 5 2 ] . In addition, concentrations of SO 2 in the flue gas a re low following the wet scrubber and there are fewer concerns related to SO 3 formation and ammonium salt deposition [ 5 2 ]

. While catalyst vendors typically provide a 24,000 hour (or 3 year) guarantee for catalysts, catalysts in tail

- end units may last for extended periods. One source c ites tail

- end SCR units in Europe that continue to operate using the initial catalyst that was installed in the 1980’s and have up to 130,000 operating hours [ 8 9 ] , and another source reports tail

- end catalysts that lasted for 100,000 operating hours [ 5 2 ]

.

Indirect Annual Costs
In general, as mentioned in the Cost Manual Methodology chapter in Section 1 of the Control Cost Manual, indirect annual costs (fixed costs) include the capital recovery cost, property taxes, insurance, administrative charges, and overhead. Capital recovery cost is basedChapter 2

– Selective Catalytic Reduction

2

- 78 on the anticipated equipment lifetime 28 and the annual interest rate emplo yed. 29 For the purposes of this cost example, t he equipment lifetime of an SCR system is assumed to be 30 years for power plants and 20 to 25 years for industrial boilers

. Th ese assumption s are based on several sources, including estimates by six petroleum refiners that SCR for fluidized catalytic cracking units and other process units would be between 20 and 30 years [22] ; results from a survey conducted by the South Coast Air Quality Management District that shows equipment life for SCRs at refineries to be 20 to 25 years, 30 an expert report in the NC law suit against TVA coal

fired electric generation units indicated e xpected useful life of an SCR is 30 years [ 90 ]; a 2002 study of the economic risks from SCR operation at the Detroit Edison Monroe power plant used 30 years as the anticipated lifetime [91 ] ; and a design lifetime of 40 years was used for an SCR at the San Juan Generating Station [9 2 ]. Thus, broadly speaking, a representative value of the equipment life for SCR at power plants can be considered as 30 years. For other sources, the equipment life c an be between 20 and 30 years. The remaining life of the boiler may also be a determining factor for the system lifetime.
In many cases

, property taxes do not apply to capital improvements such as air pollution control equipment; therefore, for this analysis, taxes are assumed to be zero [ 4 1 ]. The cost of overhead for an SCR system is also considered to be zero. An SCR system is not viewed as risk

in creasing hardware (e.g., a high

- energy device such as a boiler or a turbine). Consequently, insurance on an SCR system is on the order of a few cents per thousand dollars annually [ 4 1 ]. Finally, there are two categories of overhead, payroll and plant. Payroll overhead includes expenses related to labor employed in operat ion and maintenance of hardware, whereas plant overhead accounts for items such as plant protection, control laboratories, and parking areas. Because this procedure assumes that no additional labor is needed in operation of an SCR system, payroll overhead is zero and plant overhead is considered to be negligible.
Using these assumptions, indirect annual costs

, IDAC, in $/yr

, c onsist of both administrative charges and capital recovery, which c an be expressed as :

(2. 6 8 )

Administrative Charges
Administrative charges may be calculated as:

Administrative C h arges = 0

. 03 × ( ( Operator
Labor Cost ) + 0

. 4 × ( Annual Maintenance

Cost ) ) (2. 6 9 )

Where

28 The term “equipment life” as used here in this chapter and through t he Control Cost Manual refers to operational or design life. See Section 1, Chapter 2 for more explanation.
29 The OMB specifies an interest rate of 7% in Circular A

- 4. Regulatory Analysis. Office of Management and Budget. September 17, 2003. Availabl e at https://www.whitehouse.gov/sites/default/files/omb/assets/regulatory_matters_pdf/a

- 4.pdf
30 Presentation to NOx RECLAIM Working Group, South Coast Air Quality Management District, April 29, 2015. Presentation is available at http://www.aqmd.gov/docs/default

- source/rule

- book/Pro posed

- Rules/regxx/nox

reclaim_wgm_042915b.pdf?sfvrsn=2

.Chapter 2

– Selective Catalytic Reduction

2

- 79 Operator Labor Cost = t SCR x Operator Hours/day x Labor Rate
In general, the operating labor cost in this equation will be small because operation of an SCR system require s only minimal, operating or supervisory labor.

C apital Recovery
Capital recovery is estimated as:

CR = CRF × TCI (2. 70 )

where TCI is the total investment

, a nd CRF is the capital recovery factor and defined by:

(2. 71 )

where i is the interest rate, and n is the equipment life of the SCR system

. Total Annual Cost

The total annual cost ( TAC) for owning and operating an SCR system is the sum of direct and indirect annual costs as given in the following equation:

(2. 72 )

Cost Effectiveness
The cost in dollars per ton of NO x removed per year is:

(2. 73 )

Where:

Cost Effectiveness = the cost effectiveness, $/ton NO x Removed

/ yr = annual mass of NO x removed in the SCR, ton/yr



### 2.5 Example Problem #1

– Utility Boiler

An example problem that calculates both the design paramete rs and capital and annual
costs for an SCR system applied to a 120 MW utility boiler firing bituminous coal is presented below. The follow ing assumptions are made to perform the calculations:

Fuel High Heating Value

, HHV 1 2 ,000 Btu/lb
Net Plant Heat Rate, NPHR 10 MMBtu/MWh
Maximum Actual Output 102 MWChapter 2

– Selective Catalytic Reduction

2

- 80 Total Annual Fuel Consumption

, actual m fuel 7.45 x 10 8 lb Number of plant (boiler) operating days, t Plant 365 days Number of SCR operating days

, t SCR 36 5 days
Inlet NO Level

, NO X in 0. 35 lb/MMBtu Required Annual Average Controlled NO x Emission Level

, NO X out 0. 05 lb/MMBtu Acceptable Ammonia Slip

, Slip 2.0 ppm
Base C ase F lue G as Flow Rate Factor , Eastern Bituminous

, q fuel 484 ft 3 /min per MMBtu/hr
Fuel Sulfur Content

, S 1.0% by weight
Stoichiometric Ratio Factor for Ammonia, SFR 1.05
Stored Ammonia Concentration

, C sol 29% Number of Days of Storage for Ammonia

, t 14 days Pressure Drop for SCR Ductwork

, ∆P duct 3 inches w.g. Pressure Drop for each Catalyst Layer

, ∆ P catalyst 1 inch w.g. Temperature at SCR Inlet

, T 650°F Plant elevation , P ELEV 1,500 ft In addition to these assumptions, the estimated economic factors for the cost equations are:

Cost year 20 1 2 $ Equipment Life 3 0 years
Annual Interest Rate 7% Catalyst Cost 31 5,500 $/ m 3 ($160/ft 3 ) [ 7b ]
Electrical Power Cost 31 0.0 372 $/kWh [ 9 3 ]
29% Ammonia Solution Cost 31 0.4 7 5 $/lb of ammonia [ average of vendor quote s for 2011 and 2013 ] Operating Life of Catalyst 24,000 hours Number of hours of operator labor 4 hours/day [8] Labor Rate $60/hour (including benefits) [8] Retrofit Factor 1 (average level of difficulty)
2.5.1 Design Parameter Example #1 32

Boile r Calculations

The boiler annual heat input rate, Q B , is c alculated from the High Heating Value for bituminous coal (see Table 2.5 for typical values if the actual value is unknown) and the maximum fuel consumption rate, :

31 The electricity, catalyst, and reagent unit costs used in this example are based on data for 2012. These values are provided here for demonstration purposes only. When estimating direct annual oper ating costs, the current price of these commodities reflecting the year in which the cost estimate is made should be used. Catalyst and reagent prices can be obtained from vendors. Industrial plants should use the electricity price from their latest utilit y bill, while electricity generators should use the busbar rate.

32 Note: Results of all parameter calculations are shown rounded to an acceptable number of significant figures. However, the full, unrounded value is used in subsequent parameter and cost calculations that use the parameter as an input. Thus, the results shown for subsequent calculations often differ from what would be calculated using the shown rounded inputs. The use of extra significant figures in the subsequent calculations does not imp ly greater accuracy of the numbers.Chapter 2

– Selective Catalytic Reduction

2

- 81 The plant capacity factor is calculated from the maximum and annual MW output using Equation 2.8 a :

The SCR sy s tem capacity factor is calculated from the months of SCR operation, 12 months , using Equation 2.9 :

The total capacity factor including both plant and SCR capacity factors is given by:

CF total = 0.85 x 1 .0 = 0.85 = 85 %

The flu e gas flow rate using Equation 2. 14 is:

The NO x removal efficiency, η NOx , is calculated from the inlet NO x level and the required controlled NO x emission level using Equation 2. 10 :

SCR Reactor Calculations
The catalyst volume using Equation 2. 22 and the equations for each adjustment factor is:

Vol catalyst = 2.81 x 1, 2 00 MMBtu/hr

× [0.2869 + (1.058 x 0.86 )] (  adj x )

× [ 0.8524 + (0.3208 x 0.35 ) ] (NO x adj )

× [1.2835

– (0.0567 x 2.0)] (Slip adj )Chapter 2

– Selective Catalytic Reduction

2

- 82 × [ 0.963 6 + (0.0455 x 1.0)] (Sulfur adj )

× [ 15.16

– (0.03937 x 650) + (0.0000274 x 650 2 )] (Temperature a d j )

= 5,300 ft 3

The catalyst and SCR cross

- s ectional areas using Equations 2.2 8 and 2.2 9 are:

The length and width of the reactor using Equation 2. 30 is:

The first estimate of the number of catalyst layers using Equation 2. 31 is:

Rounding this value gives, n layer = 3.
Checking the actual catalyst height using Equation 2

. 32 :

This value is within the design height limi ts of 2.5

– 5 feet.
The total number of catalyst layers is determined by Equation 2.3 3 with one empty catalyst layer:

n total = 3 + 1 = 4

The SCR height, excluding the outlet duct and hoppers using Equation 2.3 4 is:

h SCR = 4 × ( 7 + 4.0) + 9 = 53 ftChapter 2

– Selective Catalytic Reduction

2

- 83 Reagent Calculations
The mass flow rate of the reagent is calculated using the molecular weight of the reagent, 17.03 g/mole and NO 2 , 46.01g/mole. For an SRF of 1.05, the reagent mass flow rate is given by Equation 2.3 5 :

The mass flow rate of 29% aqueous ammonia solution is given by Equation 2.3 6 :

The solution volume flow rate can then be calculated from Equation 2.3 7 where  sol is the density of the 29% aqueous ammonia soluti on, 56.0 lb/ft 3 at 60 o F

, and the conversion factor is 7.481 gal/ft 3 :

T he total volume stored in the tank( s ) is based on the volume that the SCR system requires for 14 days of operation. The onsite storage requ irement is given by Equation 2.3 8 :

The onsite storage requirement for ammonia solution is approximately 22,000 gallons per 14 days of operation. Capital Cost Elevation Factor Calculation
The elevation factor for use in calculating the SCR base unit cost and the balance of plant costs is given by Equation 2.39 with the atmospheric pressure at 1500 ft above sea level (13.9 psia):Chapter 2

– Selective Catalytic Reduction

2

- 84 2.5.2 Cost Estimation Example

Once the SCR system is sized, the capital and annual costs for the SCR system can be estimated. The TCI is estimated using Equation 2. 40 . The SCR cost , RPC, APHC and BPC must be calculated individually using equations 2.41, 2.42

, 2.43 , and 2.44

, respectively. Th ese calculations are shown below.

Annual costs are based on the economic factors listed above. In addition, the SCR system in this example is assumed to operate the entire year with a boiler loading of 85 %, resulting in a total capacity factor of 85 %. As discussed in section 2.5.1, the direct annual costs consist of the variable direct annual costs (reagent and electricity) and semivariable direct annual costs (maintenance and annual catalyst cost).

The annual maintenance costs are estimated using Equation 2. 5 7 :

An estimate for power consumption is given by Equation 2. 60 :

Electricity cost can t hen be estimated from Equation 2. 62 :

Reagent cost is estimated using Equation 2. 5 8 :Chapter 2

– Selective Catalytic Reduction

2

- 85 A portion of the catalyst is replaced every few years as discussed earlier in this chapter

. The annual catalyst replacement cost can be estimated using E quations 2.63 through 2.66 for catalyst replacement cost m ethodology 1 or using Equation 2. 6 7 for catalyst replacement cost

m ethodology 2 : Using catalyst replacement cost m ethodology 1, the actual pr ice at the time of purchase is estimated using Equation 2.63. This example assumes that one layer is replaced at a time, and based on the calculation in section 2.5.1, there are 3 catalyst layers. Thus, R layer = 3, and the total catalyst replacement cost i s given by:

To account for the time value of money, these periodic costs are amortized over the years preceding the actual catalyst purchase using the future worth factor as calculated using Equation 2.62

. The term, Y, in Equation 2. 65 is estimated using Equation 2. 66

. Assuming the boiler operates continuously all year means h year is 8,760 hr /yr, and Y is given by :

And the future worth factor for the catalyst replacement is given by:

The annual catalyst replacement cost can then be estimated using Equation 2.64 :

Alternatively, using catalyst replacement cost m ethodology 2, the annual catalyst replacement cost can be estimated using Equation 2.67:

The total direct annual cost is given by the sum of the variable direct annual costs and semivariable direct annual costs . If using cost methodology 1 for the annual catalyst replacement cost, the total direct annual cost is :Chapter 2

– Selective Catalytic Reduction

2

- 86 Alternatively, if using cost methodology 2 for the catalyst replacement cost, the total direct annual cost is:

As discussed in section 2. 4.2, property taxes and overhead are both assumed to be zero, and insurance costs are assumed t o be negligible. Thus, administrative charges and capital
recovery are the only components of indirect annual costs estimated in this analysis. Administrative charges are calculated using equation 2.6 9 as:

The capital recovery factor, CRF, is defined by Equation 2. 71 as :

and the capital recovery is calculated from Equation 2. 70 :

Capital Recovery = 0

. 0806 × $ 38

, 239

, 160

yr = $ 3

, 081

, 556

yr

The total indirect annual costs (IDAC) are calculated in Equation 2.6 8 :

The total annual cost is the sum of the direct annual and indirect annua l costs given by Equation 2. 72 . If using cost methodology 1 for the catalyst replacement cost, the total annual cost is :

Alternatively, if using cost methodology 2 for the catalyst replacement cost, the total annual cost is:Chapter 2

– Selective Catalytic Reduction

2

- 87 The annual cost in terms of NO x removed can be calculated using the total annual cost and the tons of NO x removed annually . The annual reduction in NO x emissions is given by Equation 2. 11 :

a nd the cost effectiveness is estimated using Equation 2. 73 . If using cost methodology 1 for the catalyst replacement cost, the cost effectiveness is :

Alternatively, if cost methodology 2 is used for the catalyst replacement cost, then the cost effectiveness is:



### 2.6 Example Problem #2

– Industrial Boiler

An example problem that calculates both the desi gn paramete rs and capital and annual
costs for an SCR system applied to a 550 MMBtu/hr industrial boiler firing bituminous coal is presented below. The following assumptions are made to perform the calculations:

Fuel High Heating Value, HHV 1 2 ,000 Btu/lb Maximum Fuel Consumption Rate, fuel 4

. 58 x 10 4 lb/hr
Total Annual Fuel Consumption, actual m fuel 3.30 x 10 8 lb Number of plant boiler operating days 333 days Number of SCR operating days, t SCR 333 days
Inlet NO Level, NO Xin 0. 35 lb/MMBtu Required Annual Average Controlled NO x Emission Level, NO Xout 0. 05 lb/MMBtu Acceptable Ammonia Slip, Slip 2.0 ppm
Base Case Flue Gas Flow Rate Factor , Eastern Bituminous, q fuel 484 ft 3 /min per MMBtu/hr Fuel Sulfur Content, S 1.0% by weight
Stoi chiometric Ratio Factor for Ammonia, SR F 1.05 Stored Ammonia Concentration, C sol 29% Number of Days of Storage for Ammonia, t 14 days Pressure Drop for SCR Ductwork, ∆P duct 3 inches w.g. Pressure Drop for each Catalyst Layer, ∆P catalyst 1 inch w.g. Temperature at SCR Inlet, T 650°F Plant elevation, P ELEV <500 ft above sea levelChapter 2

– Selective Catalytic Reduction

2

- 88 In addition to these assumptions, the estimated economic factors for the cost equations are:

Cost year 2012$ Equipment Life 3 0 years
Annual Interest Rate 7% Catalyst Cost 33 5,500 $/ m 3 ($160/ft 3 ) [ 7b ]
Electrical Power Cost 3 3 0.0 67 $/kWh [ 93 ]
29% Ammonia Solution Cost 33 0.4 7 5 $/lb of ammonia [ average of vendor quote s for 2011 and 2013 ] Operating Life of Catalyst 24,000 hours Number of hours of operator labor 4 hours/day Labor Rate $60/hour (including benefits) Retrofit Factor 1.4
2.6 .1 Design Parameter Example #2 34

Boiler Calculations

The boiler annual heat input rate, Q B , is c alculated from the High Heating Value for bituminous coal (see Table 2.5 for typical values if the actual value is unknown) and the maximum fuel consumption rate

, :

The plant capacity factor is calculated from the maximum and annual average fuel consumption:

The SCR system capacity factor is calculated from the fraction of boiler operating time during which the SCR also operates :

33 The electricity, catalyst, and reagent unit costs used in this example are based on data for 2012. These values are provided here for demonstration purposes only. When estimating direct annual operating costs, the curr ent price of these commodities reflecting the year in which the cost estimate is made should be used. Catalyst and reagent prices can be obtained from vendors. Industrial plants should use the electricity price from their latest utility bill, while electri city generators should use the busbar rate.
34 Note: Results of all parameter calculations are shown rounded to an acceptable number of significant figures. However, the full, unrounded value is used in subsequent parameter and cost calculations that use the parameter as an input. Thus, the results shown for subsequent calculations often differ from what would be calculated using the shown rounded inputs. The use of extra significant figures in the subsequent calculations does not imply greater accuracy of the numbers.Chapter 2

– Selective Catalytic Reduction

2

- 89 The total capacity factor including both plant and SCR capacity factors is given by:

CF total = 0.82 x 1.0 = 0.82 = 82 %

The flu e gas flow rate using Equation 2.14 is:

The NO x removal efficiency, η NOx , is calculated from the inlet NO x level and the required controlled NO x emission level using Equation 2. 10 :

SCR Reactor Calculations
The catalyst volume using Equation 2. 22 and the equations for each adjustment factor is:

Vol catalyst = 2.81 x 55 0 MMBtu/hr

× [0.2869 + (1.058 x 0.86 )] (  adj x )

× [ 0.8524 + (0.3208 x 0.35 ) ] (NO x adj )

× [1.2835

– (0.0567 x 2.0)] (Slip adj )

× [ 0.963 6 + (0.0455 x 1.0)] (Sulfur adj )

× [ 15.16

– (0.03937 x 650) + (0.0000274 x 650 2 )] (Temperature a d j )

= 2,400 ft 3

The catalyst and SCR cross

- s ectional areas using Equations 2.2 8 and 2.2 9 are:Chapter 2

– Selective Catalytic Reduction

2

- 90 The length and width of the reactor using Equation 2. 30 is:

The first estimate of the number of catalyst layers using Equation 2. 31 is:

Rounding this value gives, n layer = 3.
Checking the actual catalyst height using Equation 2

. 32 :

This value is within the design height limi ts of 2.5

– 5 feet.
The total number of catalyst layers is determined by Equation 2.3 3 with one empty catalyst layer:

n total = 3 + 1 = 4

The SCR height, excluding the outlet duct and hoppers using Equation 2.3 4 is:

h SCR = 4 × (7 + 4.0) + 9 = 53 ft

Reagent Calculations
The mass flow rate of the reagent is calculated using the molecular weight of the reagent, 17.03 g/mole and NO 2 , 46.01g/mole. For an SRF of 1.05, the reagent mass flow rate is given by Equation 2.3 5 :

The mass flow rate of 29% aqueous ammonia solution is given by Equation 2.36 :

The solution volume flow rate can then be calculated from Equatio n 2.3 7 where  sol is the density of the 29% aqueous ammonia solution, 56.0 lb/ft 3 at 60 o F, and the conversion factor is 7.481 gal/ft 3 :Chapter 2

– Selective Catalytic Reduction

2

- 91 T he total volume stored in the tank( s ) is based on the volume that the SCR system requires for 14 days of operation. The onsite storage requ irement is given by Equation 2.3 8 :

The onsite storage requirement for ammonia solution is approximately 10

, 0 00 gallons per 14 day s of operation.

Capital Cost Elevation Factor Calculation
The elevation factor for use in calculating the SCR base unit cost and the balance of plant costs is given by Equation 2.39 with the atmospheric pressure at <500 ft above sea level (14.7 psia):

2.6 .2 Cost Estimation Example #2
Once the SCR system is sized, the capital and annual costs for the SCR system can be estimated. The TCI is estimated using Equation 2.4 7 . The SCR cost , RPC, APHC and BPC must be calculated i ndividually using equations 2.4 8 , 2.4 9 , 2. 50 and 2.5 1 , respectively. These calculations are shown below.Chapter 2

– Selective Catalytic Reduction

2

- 92 Annual costs are based on the economic factors listed above. In addition, the SCR system in this example is assumed to operate all 333 days that the boiler operates, and the boiler loading is 8 2 %, resulting in a total capacity factor of 8 2 %. As discussed in section 2.6.1, the direct annual cost s consist of the variable direct annual costs (reagent and electricity) and semivariable direct annual costs (maintenance and annual catalyst cost).

The annual maintenance costs are estimated using Equation 2.5 7 :

An estimate for po wer consumption is given by Equation 2. 61 :

Electricity cost can t hen be estimated from Equation 2. 62 :

Reagent cost is estimated using Equation 2.5 8 :

A portion of the catalyst is replaced every few years as discussed earlier in this chapter. The annual catalyst replacement cost can be estimated using Equations 2.63 through 2.66 for catalyst replacement cost methodology 1 or using Equation 2.67 for catalyst replacement cost methodol ogy 2 : Using catalyst replacement cost methodology 1, the actual price at the time of purchase is estimated using Equation 2.63. This example assumes that one layer is replaced at a time, and based on the calculation in section 2.5.1, there are 3 catalyst layers. Thus, R layer = 3, and the total catalyst replacement cost is given by:Chapter 2

– Selective Catalytic Reduction

2

- 93 To account for the time value of money, these periodic costs are amortized over the years preceding the actual catalyst purchase using the future worth factor as calculated using Equation 2.62. The term, Y, in Equation 2.65 is estimated using Equation 2.66. Assuming the boiler operates continuously all year means h year is 8,760 hr /yr, and Y is given by :

And the future worth factor for the catalyst replacement is given by:

The annual catalyst replacement cost can then be estimated using Equation 2.64 :

Alternatively, using catalyst replacement cost methodology 2, the annual catalyst replacement cost can be estimated using Equation 2.67:

The total direct annual cost is given by the sum of the variable direct annual costs and semivariable direct annual costs . If using catalyst replacement cost methodology 1, the total direct annual cost is :

Alternatively, if using catal yst replacement cost methodology 2, the total direct annual cost is:

As discussed in section 2.4.2, property taxes and overhead are both assumed to be zero, and insurance costs a re assumed to be negligible. Thus, administrative char ges and capital
recovery are the only components of indirect annual costs estimated in this analysis. Administrative charges are calculated using equation 2.6 9 as:

The capital recovery factor, CRF, is defined by Equation 2. 71 as :Chapter 2

– Selective Catalytic Reduction

2

- 94 and the capital recovery is calculated from Equation 2. 70 :

The total indirect annual costs (IDAC) are calculated in Equation 2. 68 :

The total annual cost is the sum of the direct annual and indirect annua l costs given by Equation 2. 72. If using catalyst replacement cost methodology 1, the total annual cost is :

Alternatively, if using catalyst replacement cost methodology 2, the total annual cost is:

The annual cost in terms of NO x removed can be calculated using the total annual cost and the tons of NO x removed annually. The annual reduction in NO x emissions is given by Equation 2. 11 :

a nd the cost effectiveness is estimated using Equation 2. 73. If using catalyst replacement cost methodology 1, the cost effectiveness is :

Alternatively, if using catalyst replacement cost methodology 2, the cost effectiveness is:Chapter 2

– Selective Catalytic Reduction

2

- 95Chapter 2

– Selective Catalytic Reduction

2

- 96 

### References

[1] Institute of Clean Air Companies (ICAC). White Paper. Selective Catalytic Reduction
(SCR) Control of NO x Emissions From Fossil Fuel

- Fired Electric Power Plants. May

2009.

[ 2 ] Foerter

, D.

, and C. Whiteman

. Typical Installation Timelines for NO x Emissions Control Technologies on Industrial Sources

. Institute of Clean Air Companies (ICAC)

. December

2006.

[ 3 ] Salbilla, D. (Haldor Topsoe Inc.) and R. Lack (Emission Credit Brokers). Technology,
Allowances, and C redits: Components of a Sound NO x Reduction Strategy. Undated.

[4] U.S. Environmental Protection Agency. Electric Generating Unit Database. EGU_ICR_PartI_and_PartII. Based on EGU information Collection Request. December
16, 2011. Available at: http://www.epa.gov/ttn/atw/utility/utilitypg.html

.

[5] Whiteman, C. Selective Non

- Catalytic Reduction Technology Costs for Industrial Sources. Memo from Chad Whiteman, Deputy Director, Institute of Cle an Air Companies. March 18, 2009.

[6] Integer Research. Reagent Consumption Into US Coal

- Fired Power Plants to Hit 920,000 Tonnes Ammonia Equivalent in 2015. August 24, 2011. Available at: http://www.integer

research.com/2011/environment

- emissions/news/ammonia

- consumption

- coalfired

power

- plants

- hit/

.

[7 a ] Cichanowicz, J. Current Capital Cost and Cost

- Effectiveness of Power Plant Emission s Control Technologies. Prepared for UARG. January 2010.

[7b] Cichanowicz, J. Current Capital Cost and Cost

- Effectiveness of Power Plant Emissions Control Technologies. Prepared for UARG. July 2013.

[ 8 ] U.S. Environmental Protection Agency (EPA). D ocumentation for EPA Base Case
(v. 5 .1 3 ) Using the Integrated Planning Model . Office of Air and Radiation . EPAChapter 2

– Selective Catalytic Reduction

2

- 97 Publication No. #4 5 0R1 3 00 2

. November 201 3

. Available at: http://www.epa.gov/powersectormodeling/BaseCasev513.html

.

[9] Schimmoller, B. SCR Dominates NO x Compliance Plans. Power Engineering. July 2000.

[ 10 ] Hoskins, B. Uniqueness of SCR Retrofits Translates into Broad Cost Variations

. Washington Group International

. Power Engineering

. May 2003 . Available at : http://pepei.pennnet.com/display_article/176832/6/ARTCL/none/none/1/Uniqueness

- of

SCR

- Retrofits

- Translates

- into

- Broad

- Cost

- Variations/

.

[ 11 ] Khan, S., and R. Srivastava

. Up d ating Performance and Cost of NO x Control Technologies in th e Integrated Planning Model

. U.S. EPA

. Mega Symposium

. 2004.

[ 12 ] Marano , M., and G. Sharp

. Estimating SCR I nstallation C osts

. Power Magazine

. May

2006.

[ 13 ] Cichanowicz , J. Why are SCR Costs S till R ising? Power

. Vol. 148 Issue 3

. April 2004

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