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 27709

May 2016
Chapter 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 Considerations ......................................................................................... 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 NOx and Stack NOx ................................................................................... 54

2.3.5 NOx Removal Efficiency.................................................................................... 54

2.3.6 NOx 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 NOx 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............................................................................................................................... 96
Chapter 2 – Selective Catalytic Reduction

2-2

2. SELECTIVE CATALYTIC REDUCTION

2.1 Introduction

Selective catalytic reduction (SCR) has been applied to stationary source fossil fuel–fired 

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 in 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 more 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 more than 650 combined cycle gas turbines 

[1]. SCR can be applied as a stand-alone NOx control or with other technologies, including

selective non-catalytic reduction (SNCR)1 and combustion controls such as low NOx burner 

(LNB) and flue gas recirculation (FGR) [2]. 

SCR is typically implemented on stationary source combustion units requiring a higher

level of NOx reduction than achievable by selective noncatalytic reduction (SNCR) or 

combustion controls. Theoretically, SCR systems can be designed for NOx removal efficiencies 

up close to 100 percent (%). In practice, commercial 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 follows other NOx controls 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 typically achieve 

approximately 25 to 75% reduction efficiencies [5].

Either ammonia or urea may be used as the NOx 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 NOx 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 respondents 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 NOx 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 2011 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

2-3

small number of early SCR retrofits on utility boilers prior to 2000, the average costs were about 

$100/kW, in 2011 dollars, and there was little scatter in the data. From 2000 to 2007, the SCR 

costs for 32 utility boilers ranged from about $100/kW to $275/kW (2011$), 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 2011, 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 [7b]. Typical operation and maintenance costs are approximately 0.1 cents per 

kilowatt-hour (kWh) [7a, 8]. Table 2.1a 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.

The procedures for estimating costs presented in this report are based on cost data for 

SCR retrofits on existing coal-, oil-, and gas-fired boilers for electric generating units larger than 

25 MWe (approximately 250 MMBtu/hr). Thus, this report’s procedure estimates costs for 

typical retrofits of such boilers. The methodology for utility boilers also has 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 capital 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 often 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 with 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 used to develop study-level accuracy (±30%) 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 has been successfully implemented at seven European cement 

plants in Solnhofer, Germany (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, there 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 SO3 concentrations from pyritic 

sulfur found in the raw materials used by U.S. cement plants. The SO3 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 – Selective Catalytic Reduction

2-4

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. There have been concerns expressed about the potential for catalyst poisoning 

by sodium, potassium, and arsenic trioxide. Finally, other concerns 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 capital 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 – Selective Catalytic Reduction

2-5

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

Source 

Category Unit Size Fuel Type

Capital Cost

Min Avg Max $ Year Comments Reference

Electric 

Generating 

Units

NAa 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.

[7a]

~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 – Selective Catalytic Reduction

2-6

Table 2.1b: 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

IndustrialCommerci

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.

[16]

100–

1,000 

MMBtu/

hr

Oil NA ($5,550–

$11,100/MMBtu/hr)

1999$ Estimated [16]

100–

1,000 

MMBtu/

hr

Gas NA ($4,010–

$8,010/MMBtu/hr)

1999$ Estimated [16]

100 

MMBtu/

hr

Gas NA ($7,500/MMBtu/hr) 1999$b Vendor Cited source in reference [15] is an unpublished 

letter from a vendor.

[15]

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. [15]

321 

MMBtu/

hr

Wood NA ($1,980/MMBtu/hr) 2006$ Likely 

Estimated

[18]
Chapter 2 – Selective Catalytic Reduction

2-7

Source 

Category

Unit 

Size

Fuel 

Type

Capital Cost: average 

(range) $ Year

Actual, 

Vendor 

Quote, or 

Estimated? Comments Reference

Petroleum 

Refining –

Steam 

Boilers

650 

MMBtu/

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.

[19]

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. [19]

350 

MMBtu/

hr

Refinery 

oil

NA ($3,100–

$25,800/MMBtu)

2004$ c Estimated Same comment as above. [19]

10 

MMBtu/

hr

Gas or 

refinery 

fuel 

gas/NG 

combo

$19,200/MMBtu ($12,000–

$26,500/MMBtu)

1999b 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)

1999b 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)

1999b 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)

1999b 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)

1999b Vendor/

Estimated

Same comment as above. [20]
Chapter 2 – Selective Catalytic Reduction

2-8

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. [15]

Petroleum 

Refining –

FCCU

70,000 

barrels/s

tream

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)

1.09

million 

short tpy 

clinker

NA NA ($6.9 per short ton

clinker)

2006a 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)

2006a Estimated Same comment as above. [24,25]

2.16 

million 

short tpy 

clinker

NA NA ($3.9 per short ton 

clinker)

2006a Estimated Same comment as above. [24,25]
Chapter 2 – Selective Catalytic Reduction

2-9

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)

2011 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)

2006a 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)

2006a 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. [15]

80 MW Gas NA ($51/kW) 1999$a Vendor Retrofit cost, excluding balance of plant costs. [15]

2 MW Gas NA ($237/kW) 1999$a Vendor Retrofit cost. [15]

12 MW Gas NA ($167/kW) 1999$a Vendor Retrofit cost. [15]

Internal 

Combusti

on Engine

1,800 hp Diesel 

(No. 2 

fuel oil)

NA ($0.18 million) 1994 Actual New cost [15]

a Year of reference.

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

c Commission year of the SCR.
Chapter 2 – Selective Catalytic Reduction

2-10

2.2 Process Description

Like SNCR, the SCR process is based on the chemical reduction of the NOx 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, ammonia 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 NOx within a specific temperature range and in the presence of the catalyst and oxygen 

to reduce the NOx into molecular nitrogen (N2) and water vapor (H2O).

The use of a catalyst results in two primary advantages of the SCR process over SNCR. 

The main advantage is the higher NOx 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 operating 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 replacement cost has been the largest cost, the reagent 

cost has become the most substantial portion of operating costs for most SCR [7b].

4

Figure 2.1 shows a simplified process flow schematic for SCR. Reagent is injected into 

the flue gas downstream of the combustion 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, NOx 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 combustion unit where SCR 

systems can be installed. 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 particulate 

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) is 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/m3) [30].
Chapter 2 – Selective Catalytic Reduction

2-11

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 

(NH3) because it readily penetrates the catalyst pores. The ammonia, either in anhydrous or 

aqueous form, is vaporized before injection by a vaporizer. Within the appropriate temperature 

range, the gas-phase ammonia then decomposes into free radicals, including NH3 and NH2. After 

a series of reactions, the ammonia radicals come into contact with the NOx and reduce it to N2

and H2O. Since NOx includes both NO and NO2, the overall reactions with ammonia are as 

follows:

NO NH O N H O

catalyst

3 2 2 2 3 2

2

1

2  2   

(2.1a)

NO NH O N H O

catalyst 2 2  4 3  2 3 2  6 2

 (2.1b)

The equations indicate that one mole of NH3 is required to remove one mole of NO and two 

moles of NH3 are required to remove one mole of NO2. However, Equation 2.1a is the 
Chapter 2 – Selective Catalytic Reduction

2-12

predominant reaction because 90 to 95% of NOx in flue gas from combustion units is NO. 

Hence, about one mole of NH3 is required to remove one mole of NOx. 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 ammonia and gas-phase NO to 

form an activated complex. The catalytic reaction, represented by Equations 2.1a and 2.1b, 

results 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 to free radicals and provides 

the activation energy for the reaction. The reaction also requires excess oxygen, typically 2–4%, 

to achieve completion. NOx reduction with ammonia is exothermic, resulting in the release of 

heat. However, because the NOx concentration 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 NOx reduction 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-ondemand”). Anhydrous ammonia is nearly 100% pure ammonia. It is a gas at normal atmospheric 

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 

concentration of 29.4% ammonia in water, although some applications use a 19% solution [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 ammonia and also requires 

shipping costs for the water solvent in the solution. Although the 29.4% solution has substantial 

vapor pressure at normal air temperatures, a vaporizer is generally required to provide sufficient 

ammonia 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 equipment. 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 NOx removed is higher for urea due to urea’s much higher molecular weight. For 

example, one SCR supplier estimated capital costs for a 130 lb/hr ammonia system to be 

$280,000 for anhydrous ammonia, $402,000 for 19% aqueous ammonia, and $750,000 for urea

[34]. Another reference reported that the equipment cost for urea is generally twice the 
Chapter 2 – Selective Catalytic Reduction

2-13

equipment 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, one reference estimated annual operating costs for 19% 

aqueous ammonia are 50% higher, 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 operating costs for anhydrous 

ammonia [35]. One reference estimated energy costs for an unspecified application to be 

$167,000 for a urea system, $73,000 to $117,000 for aqueous 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 examples in section 2.5 illustrate the procedures 

for a system using 29% aqueous ammonia as the reagent). 

Table 2.2: Ammonia 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 NH3)

Molecular weight of reagent 17.03 17.03 (as NH3)

Ratio of ammonia to solution 99.5% (by weight of NH3) 29.4% (by weight of NH3)

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% NH3 (by volume) 16–25% NH3 (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 copperbased 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 occurs, 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 proprietary. 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 (Ti), and tungsten 

(W), which significantly reduced catalyst cost. In the 1980s, metal oxides such as titanium oxide 

(TiO2), zirconium oxide (ZrO2), vanadium pentoxide (V2O5), and silicon oxide (SiO2) were 

employed to broaden the reaction temperature range. Zeolites, crystalline alumina silicates, were 

also introduced for high temperature (675–1000°F; 360–540°C) applications; however, zeolites 
Chapter 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/m3to a range of $5,000 to $6,000/m3(costs are in 2011$) [7b].

5 This 

reference also reported that catalyst prices remained in the approximate range of $5,000 to 

$6,000/m3through 2012.

Improvements to the catalyst formulations over time have decreased unwanted side 

reactions such as sulfur oxide conversions (SO2 to SO3) and increased 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 electrostatic precipitators [ESPs]) [40]. Improved 

catalyst designs have also increased catalyst activity, surface area per unit volume, and the 

temperature range for the reduction reaction. As a consequence, there is a corresponding 

decrease in the required catalyst volumes and an increase in the catalyst operating life. For coalfired boiler applications, SCR catalyst vendors typically guarantee the catalyst for an operating 

life ranging from 8,000 to 24,000 hours [1]. Applications using oil and natural gas have a longer 

operating life, typically greater than 32,000 hours [41]. In addition, operating experience 

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 NOx removal of 90%; control of residual NH3 to 2 ppm (i.e., ammonia slip); guarantees 

for SO2 oxidation to less than 1%, and in many cases, to less than 0.5%; being able to withstand 

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 contain additional compounds or supports to give 

thermal and structural stability or to increase surface area [42]. Catalyst configurations are 

generally ceramic honeycomb and pleated metal plate (monolith) designs in a fixed-bed reactor, 

which provide high surface area to volume ratio. Pellet catalysts in fluidized beds are also 

available. Pellets have greater surface area than honeycombs or pleated plates but are more 

susceptible to plugging. This limits the use of pellets to clean-burning fuels such as natural gas.

Catalyst elements 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 the large catalyst 

modules into the reactor from either the interior or exterior of the reactor, depending on the 

reactor design.

Catalysts greatly accelerate the NOx reduction reaction rate, but some catalysts have more 

favorable properties for a given application. Performance requirements that drive the choice of 

catalyst include reaction temperature range, flue gas flow rate, fuel source, catalyst activity and 

selectivity, SO2 oxidation, and catalyst operating life. In addition, the design must consider the 

cost of the catalyst, including disposal costs [33]. In the past, the initial charge of catalyst costs 

accounted for 20% or more of the capital costs for an SCR system [33], however, as catalyst unit 

cost has declined over time, this 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 from 1980 to 2006, the cost of catalyst dropped by 75% from approximately

$16,000/m3to less than $4,000/m3[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 the catalyst is reactivated (i.e., rejuvenated or regenerated) for reuse or 

its components are recycled for other uses [42]. 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 [42].

Catalyst cleaning typically means the removal of physical restrictions to 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 [43]. These activities may occur online/in-situ or offline, and may occur onsite or 

offsite [43]. The catalyst layers may be removed and transported to the cleaning, rejuvenation, or 

regeneration 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) [44].

Water-based cleaning can also be conducted to remove physical materials that plug or blind the 

catalyst [45].

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

some applications, can be considered equivalent to new catalyst [46]. 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 [43]. Today, 

regeneration processes focus on increasing the longevity 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 remove 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 

catalyst [47]. Because regeneration may cause catalysts to lose mechanical strength, 

recalcination may also be conducted to ensure catalyst impregnation and that catalyst mechanical 

strength is regained [43, 45, 47]. Improvements to the catalyst activity relative to certain 

reactions can be made using a regeneration process (e.g., impregnating other chemicals in the 

catalyst to reduce the conversion of SO2 to SO3, or to increase the oxidation of Hg over the 

conversion rate of the original catalyst while maintaining the same catalyst activity for NOx)

[47]. 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 [44].

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]; same 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 SO2 oxidation for some catalysts (i.e., SO2/SO3

conversions is no higher than the new catalyst guarantee) [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 regenerated catalysts reduces catalyst replacement cost and minimizes the need to 

dispose of spent catalyst [48]. It is estimated that a typical 500-MW coal-fired power plant will 
Chapter 2 – Selective Catalytic Reduction

2-16

spend approximately $2,000,0006 on a single layer of new catalyst [48]. The cost for regenerated 

catalyst for this same facility would be approximately $1,000,000 for a single layer of catalyst

[48]. Disposal costs when replacing a spent catalyst could be $50,000–$200,000 per layer, and 

these costs are avoided with regenerated catalysts [48]. 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 NOx removed from the flue 

gas. The major design and operational factors that affect the NOx removal performance of 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 and the combustion gases;

 Molar ratio of injected reagent to inlet NOx;

 Inlet NOx 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 operational 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;

 SO2 and SO3 concentrations in gas stream;

 Catalyst deactivation; and

 Catalyst management.

The major differences between SNCR and SCR are discussed below.

Temperature

The NOx 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 NOx

reduction 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 [Reference 48].
Chapter 2 – Selective Catalytic Reduction

2-17

(N2O) formation. At temperatures above the specified range, N2O formation 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 catalysts (metal 

oxides), the operating temperatures for the SCR process range from 480°F to 800°F (250–430°C) 

[46]. Figure 2.2 is a graph of the NOx removal efficiency as a function of temperature for a 

typical metal oxide catalyst [46]. The figure shows that the rate of NOx removal increases with 

temperature up to a maximum between 700°F and 750°F (370–400°C). As the temperature 

increases above 750°F (400°C), the reaction rate and resulting NOx removal efficiency begin to 

decrease.

Figure 2.2: NOx Removal versus Temperature [46]

As flue gas temperature approaches the optimum, the reaction rate increases and less 

catalyst volume achieves the same NOx 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 

range, 700–750°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 temperature, catalyst volume, and NOx removal are 

complicated functions of the catalyst formulation and configuration. The physical 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. Therefore, the selection of the catalyst is critical 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 metal oxide–based catalysts. Reheating of the flue 

gas may be required for reactors located downstream of the air preheater. Reheating significantly 

increases SCR operational costs. This continues to be true despite natural gas prices that are 

relatively low 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 sft3[52]. If the natural gas unit cost were $5/1,000 sft3, 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 systems tolerate 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% load, 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 

with the relatively cooler flue gas exiting the economizer. An economizer feedwater bypass also 

raises the flue gas temperature. The use of an economizer bypass results in less energy transfer to 

the feedwater for steam generation; consequently, there is a small reduction in boiler efficiency. 

Lower boiler efficiencies require more fuel to be burned to meet the required boiler steam output.

Residence Time and Space Velocity

Residence time is the time the reactants are within the reactor. Higher residence times 

generally result in higher NOx 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 NOx

removal efficiency increases with decreasing 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 those 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 pore surface area (specific surface area). For coal-fired 

boilers, typical specific surface areas range from 90 to 3,800 square feet per cubic feet (ft2/ft3) 

(300–1,200 square meters per cubic meters [m2/m3]) [33]. Increasing the catalyst specific surface 

area increases the NOx removal for a given flue gas flow rate. This can be accomplished by 

either increasing the catalyst volume, which increases the reactor size, or increasing the pore 

space of the catalyst, which generally increases the catalyst cost.

Degree of Mixing

The reagent must be dispersed and mixed throughout the flue gas to ensure sufficient 

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 reagent. Some systems inject the ammonia with a carrier 

fluid, such as steam or air, to increase penetration into the flue gas. Injection systems are 

application specific. Numeric modeling of the flue gas and reagent flow optimizes the design of 

the injection system (see Section 2.2.6).

Mixing of the flue gas and ammonia occurs before entering the SCR reactor. If mixing is 

not adequate, the NOx reduction is inefficient. SCR designs must incorporate adequate duct 
Chapter 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;

 Increasing the energy imparted to the injected fluids;

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

 Modifying 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 NOx removed) 

defines the quantity of reagent needed to achieve the targeted NOx 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 NO2 can be removed with two moles of ammonia. Since NOx is 

mostly comprised of NO (approximately 95%), the theoretical stoichiometric ratio for NOx is 

close to 1.05 mole of ammonia per mole of NOx. Hence, SCR systems typically employ a 

stoichiometric ratio of 1.05 moles of ammonia per mole of NOx [33]. This assumption of an 

almost one-to-one linear relationship between the quantity of reagent and the NOx removed is 

good up to about 85% NOx reduction [46]. Above 85%, the removal efficiency begins to level 

off and more than the theoretical amount of ammonia is required for additional NOx removal

because of reaction rate limitations. Because capital and operating costs depend on the quantity 

of reagent consumed, the stoichiometric ratio factor is an important design parameter that is 

determined by the SCR designer.

Inlet NOx Concentration

The concentration of the reactants also affects the reaction rate of the NOx reduction

process. In general, higher inlet NOx inlet concentrations result in higher NOx removal 

efficiencies due to reaction kinetics [33]. However, NOx concentrations higher than 

approximately 150 parts per million (ppm) generally do not result in increased performance. Low 

NOx inlet levels result in decreased NOx 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 controls 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 NOx levels, such as natural gas–fired 

boilers.

For a given NOx removal efficiency, higher NOx levels at the SCR inlet require more

catalyst volume. For example, to achieve 90% NOx removal requires 10% more catalyst at an 

inlet NOx level of 1.7 lb/MMBtu versus an inlet NOx level of 0.8 lb/MMBtu [53]. 

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 discussed in Chapter 1, SNCR, including 

health effects, visibility of the stack effluent, salability of the fly ash, and formation of 
Chapter 2 – Selective Catalytic Reduction

2-21

ammonium sulfates. Limits 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 designed SCR systems, which operate close to the 

theoretical stoichiometry and supply adequate catalyst volume, maintain low ammonia 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 increases 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 and 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 [54]. Another method for quantifying ammonia slip is to determine the 

ammonia concentration in collected fly ash [55].

Raw materials at some cement kilns contain constituents that release ammonia to the kiln 

gas stream when heated to high temperature. Some 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 NOx reduction

reaction rate. A higher catalyst activity results in a faster reaction rate and more NOx removal. 

Catalyst activity is a function of many variables, including catalyst composition and structure, 

diffusion rates, mass transfer rates, gas temperature, and gas composition [56]. As the catalyst 

activity decreases, the NOx reduction reaction rate also decreases. This results in lower NOx

removal and higher ammonia slip levels.

The following equation describes the deactivation of the catalyst activity, K, with 

time, t [57]:

( t / )

o K K e





(2.2)

where Ko is the original catalyst activity and  is the catalyst operating life time constant. 

Figure 2.4 shows a typical catalyst deactivation curve based on Equation 2.2. As the catalyst 

activity decreases, the NOx 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 

regenerated, or new catalyst must be installed.

Catalyst Reaction Selectivity

SCR favors the NOx 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 (SO3) and nitrous oxide (N2O). SO3 is formed by the oxidation of SO2 to SO3, 

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 materials is oxidized and released into the kiln gas stream [94]. Sulfur 

oxides (SOx) are regulated under the 1990 Clean Air Act. SO3 reacts with ammonia in the flue 

gas to form ammonia sulfates. Ammonium sulfur salts deposit on the catalyst and on 

downstream equipment such as the air preheaters. SO3 and ammonium 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 percent or the SO2

content of the fuel is greater than or equal to 3 lb/mmBtu. Increasing the number of catalyst 

layers, while increasing NOx removal efficiency, can also lead to an increase in the conversion of 

SO2 to SO3.

9 In addition, SO3 emissions cause “blue plume” from the unit and have become an 

emissions concern. Newer catalysts, however, have been developed that limit the formation of 

SO3 [37], and these catalysts are now commonly used in SCR installations in the U.S. N2O is 

both an ozone depleter and a greenhouse gas. N2O has a global warming potential (GWP) of 298

as compared to carbon dioxide (CO2).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 N2O 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 at 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 KO = 24.12;  = 55,000

Pressure Loss

The flue gas pressure decreases 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 the 

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 installed. 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 kilns can be in excess of 80 g/Nm3[94]. In addition, cement dust can be sticky, creating 

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 additional catalyst cleaning and 

catalysts with larger pitch [94].

Ash Management and Dust Loading

Ash 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 amount 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 reduction 
Chapter 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/m3) [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 ensure that ash 

will not deposit and bridge over catalyst cells and pores. Plugging of the catalyst reduces the 

effective surface area by decreasing the number of active sites available for the NOx reduction 

reaction.

For high-dust applications, such as cement kilns and coal-fired boilers, catalysts with 

larger openings or pitch should be used to reduce the potential for catalyst plugging. However, 

the wider pitch reduces the surface area per unit volume, and may necessitate additional catalyst 

volume to meet emissions limits for NOx and ammonia slip.

Figure 2.5: Pitch for a Honeycomb Catalyst Configuration

SO2 and SO3 Concentrations in Gas Streams

Sulfur in the fuel forms SO2 during combustion in the boiler, i.e., boiler conversion rate, 

and some of the SO2 is further oxidized to SO3 within the boiler. In a coal-fired boiler, 

approximately 0.5 to 1.5 percent of SO2 is converted to SO3 at typical combustion temperatures, 

however, as the temperature drops more SO3 can be formed.11 Additional SO3 is formed when 

SO2 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). SO3 will combine with water to form sulfuric 

acid (H2SO4) or sulfuric acid mist. Concentrations of SO3 and H2SO4 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 SO3 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, sodium, arsenic, 

chlorine, fluorine, and lead. These constituents deactivate the catalyst by diffusing into active 

pore sites and occupying them. Catalyst poisoning represents 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 stream 

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 poison 

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 CaSO4 in cement kilns, which can reduce catalyst activity through 

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 

deposits on the surface and in the active pore sites of the catalyst. This results in a decrease of the 

number of sites available for NOx 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 edges or increased structural strength are less 
Chapter 2 – Selective Catalytic Reduction

2-26

susceptible to erosion. However, increasing catalyst 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 different physical and chemical 

properties. 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 optimal 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 

and 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 basis, 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 another option that prevents 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 particles, 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 activity 

decreases with time, the NOx 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 is the time the catalyst 

activity for a given catalyst volume (layer volume) maintains ammonia slip below the design 
Chapter 2 – Selective Catalytic Reduction

2-27

limit. Currently, vendor-guaranteed 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 [58]). In the past, CMP descriptions 

focused mainly on the catalyst replacement 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 NOx SIP Call (1998) requirements typically called for ozone season 

operation only, year round operation is necessary 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 NOx credits, or to comply with settlement agreements with the U.S. EPA and 

Department of Justice. Continuous, ongoing collection and documentation of data on plant 

loading and cycling, fuel demands and variation, and ongoing NOx performance and SO2

conversion, which can then be compared to catalyst activity data, is conducted to create the plant 

operating history. Some companies have developed computer software that collects these data 

and optimizes the costs for catalyst replacement options. In general, an annual SCR system 

inspection is conducted on the catalyst, the reactor, and the complete NH3 storage and injection 

system. Inspection of the catalyst includes a physical inspection along with catalyst sampling and

analysis on a bench-scale reactor for activity, SO2 to SO3 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/velocity and uniform NH3/NOx molar distribution. Poor 

distribution of the NH3/NOx decreases the NOx reduction and increases the NH3 slip [59]. In situ 

measurements 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 

NH3/NOx ratio and measurement of inlet and outlet NOx 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 [61]

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 that 

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, part 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 are needed, 

review of recent catalyst technology advances for newer catalysts that achieve mercury 

reductions, lower SO2 conversion rates, and lower load and temperature operation is advised, 

although some regeneration processes may offer improvements with these catalyst capabilities as 

well. Typically, the 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 [59]. There would likely be additional costs 
Chapter 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, catalyst cost was a significant portion of the annual cost of operating an SCR 

system. Under the latest operating approaches that involve using a CMP, only a fraction of the 

total catalyst inventory, rather than the entire volume, is replaced at any one time. This 

distributes the catalyst replacement costs more evenly over the lifetime of the system, and use of 

regenerated catalyst may also reduce the overall annual costs [59]. 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% tailend [62].

12 More recently for the U.S, it was reported 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. [52]. Highdust is generally considered the most economical and straightforward design provided sufficient 

space is available to construct the SCR close 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 simple 

cycle. The various configurations for boilers and gas-fired turbines are discussed below. In 

addition, there are two 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 system 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 optimal

temperature window for NOx 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 

downward through the catalyst. The reactor generally contains multiple layers of catalyst. The 

volume of catalyst required varies with each installation, as discussed previously. Soot blowers 

or sonic horns are installed to remove particulates from the catalyst surfaces. For designs that use

a honeycomb catalyst, the catalyst 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 blowers 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 [52].

 

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 is 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 SO2 (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 [42]

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. Figure 2.8 shows a low dust configuration, which locates the 

SCR reactor downstream of the ESP. In this location, the flue gas is relatively dust free. 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 poisons in the SCR reactor. In addition, low-dust SCR configurations do 

not need ash hoppers. For designs employing honeycomb catalyst, the catalyst pitch can be 

reduced to approximately 4–7 mm, resulting in lower catalyst volume. Low-dust SCR typically 
Chapter 2 – Selective Catalytic Reduction

2-31

requires only 2 layers of catalyst [52]. Longer catalyst life, lower catalyst volume, and the 

elimination of the ash hopper mean lower costs for low-dust SCR compared to high-dust 

configurations. The only disadvantage of low-dust SCR is the temperature drop of 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 temperature 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 [42]

Tail-End 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 catalysts before 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]. 

Because the flue gas temperature at the tail end is below the range required for the 

NH3/NOx 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 recuperating gas-to-gas heater. Some use catalysts specially designed for 
Chapter 2 – Selective Catalytic Reduction

2-32

temperatures between 300 – 550oF and low pressure drops, which lowers the costs flue gas 

reheating [103, 106, 107]. 

A tail-end system may have higher capital and operating costs 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 

[52], 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 SO2 in the flue 

gas for tail-end units, the catalyst lifetime is significantly increased and less expensive catalyst 

may be used [52]. Some sources have reported 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 SO3. 

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]. They 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 with 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 combustion 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 NOx 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 NOx 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 sft3 natural gas) [52]. For 

comparison, the annual reheating cost for natural gas burner would be $7.8 million/yr (assuming approximately 

$5/1000 sft3 natural gas).
Chapter 2 – Selective Catalytic Reduction

2-33

Figure 2.9: Tail-end SCR Arrangement [42]

Gas Turbines

Natural gas–fired turbine applications frequently use SCR technology for postcombustion NOx 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 within 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-temperature 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 [42].
Chapter 2 – Selective Catalytic Reduction

2-34

Figure 2.10: SCR Arrangement for a Combined-Cycle Gas Turbine [42]

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 kilns require catalyst cleaning mechanisms. The “low dust” and “tail-end” configurations 

avoid the costs of catalyst cleaning systems. Currently, three “high-dust” SCR systems14, three 

“low-dust” SCRs15, and one “tail-end” SCR16 are known. The “high-dust” SCRs reportedly 

achieve control efficiencies of approximately 80% with inlet dust loading of up to 100 g/m3. The 

“low-dust” SCRs are reported to have dust loadings less than 20 mg/m3, while the inlet dust 

loading for the “tail-end” SCR is reported to be less than 10 g/m3[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 of SCR reactors: full SCR and in-duct SCR. Full SCR 

designs house the catalyst in a separate reactor chamber. The boiler flue gas must be ducted 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 NOx 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 preheater 

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 Sarchi SCR is installed on a small Polysius Lepol kiln with no 

particulate controls, but low dust loading (reportedly less than 15 g/m3).

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 ductwork often necessitates upgrades to the 

draft fan system.

In-duct (inline) 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 NOx control technologies [41]. 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 boilers frequently employ full SCR reactors but 

may apply in-duct SCR reactors where space limitations restrict the installation of a full reactor 

[41]. Cement kilns have also used full scale SCR reactors, rather than in-duct SCRs. The SCRs 

used for cement kilns 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 layers, 

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 DeNOx®, an ammonia-based NOx reduction 

system developed and patented by Exxon Research and Engineering Company in 1975. An SCR 

system consists of five basic steps:

 Receive and store the ammonia (or the urea reactant, followed by onsite ammonia 

production);

 Vaporize 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 reduce the NOx.

Although the basic steps in an SCR system are similar for all configurations, the system 

design and equipment specifications are somewhat different. A discussion 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 example, 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. Simplified system flow schematics 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 ft3 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 steel 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 steel, 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

Economizer 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 fly 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, NOx emissions, etc., 

including NOx 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 several reagents may be used in an SCR system, 

including anhydrous ammonia, aqueous ammonia, or urea. In the past, reagents have 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 survey in 2009, about 80 percent used 
Chapter 2 – Selective Catalytic Reduction

2-37

ammonia, and 20 percent used urea [4]. Urea reagent is mostly used in SNCR systems [65], 

however, U.S. cement plants typically use 19% aqueous ammonia for SNCR systems and likely 

would use the same reagent for SCR applications. Another 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 system 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 vaporize or decompose 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 [65]. 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 NH3 [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 NH3, with a total capital investment of $175 million for the SCR system (not including the 

ammonia system) [66].

Table 2.4: Comparison of Ammonia Delivery Systems [65]

Measure Anhydrous NH3 Aqueous 19% NH3 Aqueous 29% NH3 Urea-derived NH3

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 Lowest 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 (29%) and anhydrous ammonia.

If facilities receive anhydrous or aqueous ammonia from offsite, it is received via a tanktruck or rail car and pumped into one or more storage tanks. Ammonia is typically stored as a 

liquid in horizontal cylindrical tanks. An aqueous 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 are 
Chapter 2 – Selective Catalytic Reduction

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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 

ranging from 10,000 to 20,000 gallons per tank to maintain sufficient volume 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 the 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 storage 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 ratio of about 20:1 (air:NH3), 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

Use of onsite ammonia production systems for feed into SCR reactors has increased, 

mostly as a result of safety concerns [37]. Several process types can produce ammonia from 

urea, including (1) systems that hydrolyze an aqueous urea solution to form ammonia and carbon 

dioxide (and water); (2) systems 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) systems that atomize an aqueous urea solution in a decomposition chamber 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 [65]. Almost all urea-to-ammonia 

systems use solid urea [65]. Urea-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 [65]. 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, the 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, dry urea is fed batch-wise from delivery trucks directly to the dissolver along with 
Chapter 2 – Selective Catalytic Reduction

2-39

deionized water, and the urea solution is stored. In the solution feed and control system, urea 

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 310°F (140–

150°C). The urea-to-ammonia reaction occurs in two steps: the first reaction produces 

ammonium carbamate (NH4COONH2) 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

[67].

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 injectors. The decomposition occurs under a specific temperature and 

residence time, with the decomposition temperature ranging from 600 to 1000°F (320–540°C), 

and the urea is decomposed to ammonia and isocyanic acid. The outlet ammonia stream from the 

decomposition chamber feeds into the AIG system for the SCR [68].

Figure 2-11: Urea-Derived Ammonia Production System Using U2A system [65]

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 NOx ULTRA System [68]

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 flowbalancing skid consists of flow meters and manual valves to adjust the flow to each part of the 

AIG.

The AIG consists of a network of pipes or lances connected in parallel and perforated 

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 

the flue gas. The spray angle and velocity of the injection control the trajectory of the ammonia. 

Injectors are subject to high temperatures and flue gas impingement, which cause erosion, 

corrosion, and degradation of structural integrity. Therefore, injectors are generally constructed 

of stainless steel and designed to be replaceable. Multiple injection 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 or 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 

typically use high-energy systems. High-energy systems are more expensive 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 modeling and numerical flow modeling are generally performed for 

the AIG and SCR system to ensure uniform mixing and dispersion before the gases enter the 

SCR reactor. If duct length is inadequate to ensure thorough mixing or results from the model 

study indicate 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 NOx, and inlet gas temperatures set the feed-forward signal to establish the base 

ammonia injection rate. A feedback signal measuring the SCR outlet NOx concentration is used 

to trim the base ammonia injection rate.

Although not necessarily required to achieve high NOx removal efficiencies and low 

ammonia slip, most SCR installations today employ some type of static mixer to achieve good

NH3/NO mixing [37]. Static mixers provide a more uniform flux of NOx and more uniform 

temperature and mixing of NH3 and NO [37]. For example, achieving 90% NOx removal and a 2-

ppm ammonia slip typically requires NH3/NO uniformity less 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]. The costs for static mixers vary (e.g., a 500 MW unit may have an installed cost of 

$750,000 and an additional 1 in. w.g. of flue gas pressure drop) [37]. One specific type of static 

mixer is the Delta WingTM 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 ammonia injection nozzles are located in the vortex zone immediately downstream of 

the mixer [69]. 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 determine 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 approach that may help retain good NH3/NO uniformity is to conduct an annual 

tuning of the AIG, which can improve the NH3/NO mixing (reduces NH3/NO “unmixedness” by 

2–5% on a root mean square basis) [37]. Data have shown 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 nearstartup mixing uniformity [37]. Annual AIG tuning can cost from $30,000 to $50,000 depending 

on the unit size [37]. Depending on the type of mixer used, annual tuning may not be necessary, 

because some 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 

commissioning and for annual maintenance [71].

Use of static mixers and annual tuning can either increase the NOx efficiency at the same 

ammonia slip level or extend the catalyst life at the same NOx removal efficiency [37].

Extending the catalyst life can significantly reduce operating costs, even after accounting for the 

outsourced tuning costs [37].
Chapter 2 – Selective Catalytic Reduction

2-42

Catalytic Reduction of NOx

The catalytic reduction of NOx in the SCR reactor occurs when the NOx and ammonia 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 supplier 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 Ductwork

In retrofit installations, new ductwork is required to integrate the SCR system with the 

existing equipment. In high-dust SCR systems for utility and industrial boilers, the reactor is 

located between the economizer outlet and the air heater inlet. In low-dust SCR systems for 

utility and industrial boilers, 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 the 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 zeroleakage 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 NOx 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 gases 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 

shutdown that could preclude the need for a bypass, however, particularly for SCRs that operate 

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 may 
Chapter 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 SCR bypasses, according to a report by a major engineering 

services firm.19

Soot Blower or Sonic Horn

In coal-fired boilers, soot blowers 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 blowers also keep the air heater gas passages open and thereby 

reduce system pressure drop. This is especially true for SCR retrofits where the air heater plate 

spacing is generally narrow, making it more susceptible to fouling or clogging by ammoniasulfur salts.

Retractable rake-type soot blowers that use steam or air for blowing are used in SCR 

designs. The soot blowers 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 hours, but is usually done infrequently. In European SCR 

installations, soot blowing is done approximately once or twice a week [72]. 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 blowers [37]. Sonic horns require 

less preventive maintenance than soot blowers, 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 blowers [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 blowers [37]. Although sonic 

horns may have some advantages over soot blowers, the demand for high NOx removal 

efficiency requires extremely clean catalyst and thus the best cleaning system (regardless of 

cost); therefore, higher costs for soot blowers 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 industrial 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 NOx

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 Successful 

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/brochures/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 or 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 the 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 has 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 [73].

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; and (4) use active cleaning systems. The open flow area and velocity 

recommendations 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 screens 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 erosion-tolerant 

design for high flue-gas velocity and exotic construction material [37]. The installed cost of 

screens for two SCRs on twin boilers that collectively generate 745 MW was $600,000 in 2004 

[74]. A redesign and replacement of this screen along with CFD flow modeling in 2009 cost 

$806,000 [75, 76]. Operating costs 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 SO3 formation, and also minimizes slag and fouling. 

This slag 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 been 
Chapter 2 – Selective Catalytic Reduction

2-45

demonstrated for more than six years. After successfully implementing the targeted in-furnace

injection, the facility was able to remove their pre-existing LPA screens [77].

Economizer Bypass Duct

Although the SCR reaction occurs within a temperature window of 600–750°F (320–

400°C), the catalyst for a given application is designed 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 NOx reduction reaction. 

When the economizer outlet 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 mixed 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 

gas 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) Fan

The new ductwork and the SCR reactor’s catalyst layers decrease the flue gas pressure. 

To 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. Replacement 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 Equations 2.57 and 2.58 for estimates of the total additional 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 industrial boilers use coal, distillate oil, residual oil, natural gas, and a variety 

of other fuels such as biomass (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, which 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 [53].

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 ash, 
Chapter 2 – Selective Catalytic Reduction

2-46

SO2, SO3, arsenic, and other trace pollutants than does flue gas from burning oil or natural gas. 

Natural gas is the cleanest fuel and contains the least nitrogen; therefore, burning natural gas 

results in the least NOx and pollutants in the flue gas stream. 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 SOx

Sulfur trioxide (SO3) forms during the combustion of fuels that contain sulfur, and 

additional SO3 is formed over the SCR catalyst. SO3 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 SO2 and SO3 in the gas stream. SO3 reacts with ammonia in the flue gas 

downstream of the reactor (ammonia slip) to form ammonium bisulfate (NH4HSO4) and 

ammonium sulfate [(NH4)2SO4]. 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 imposed as part of the SCR design requirements to avoid impacts on 

downstream equipment.

There are several methods for limiting the impact of ammonia-sulfur salt deposition. 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 plant. 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 SO3 concentrations raise the acid dew point of the flue gas. This 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 Ljungström© 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 salability of the ash as a byproduct 

and the storage and disposal of the ash by landfill [51].

Oxidation of SO2 in the flue gas stream to SO3 over the catalyst bed creates even more 

SO3 (a PM2.5 precursor compound) and is an 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 than 2% sulfur content or higher.

20

 In some areas of the country, 

limits for SO3 of 5 ppm or less have been implemented to control SO3 plumes emitted from SCRs

(i.e., blue plumes) [37]. Mitigation of SO3 formation from oxidation of SO2 on the catalyst has 

evolved to be of equal importance to NOx control in SCR design [37]. Catalysts designed for low 

SO2 oxidation have been developed, but depending on the catalyst supplier, their use may

increase the catalyst volume needed to meet the same NOx removal and ammonia slip 

 

20 U.S. EPA. Regional Haze FIP for New Mexico (San Juan NOx 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 

mitigating SO3 formation that has been analyzed is use of reagents such as sodium, magnesium, 

or calcium-based sorbents [37]. However, the use of reagents alone to lower SO3 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 corrosion and fouling of ductwork, ESPs, and the blade housings

of ID fans; eliminating modification or coatings for the air heater; and reducing the heat rate and 

fuel costs [37, 78]. Use of low-SO2-oxidation catalyst can be used in combination with reagentbased SO3 mitigation to lower costs of reagent-only approaches to reducing SO3 [37].

Formation of Arsenic Oxide

Arsenic oxides (As2O3) formed during combustion of fuel containing 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 physically block active sites by condensing in the pores, and 

As2O3 can react with oxygen on the vanadium compounds in the catalyst to form As2O5, which 

chemically bonds to the site [79]. Catalyst manufacturers have developed arsenic-resistant 

catalysts that include control on porosity to minimize pore condensation and addition of 

compounds such as molybdenum or other materials to mitigate deactivation of active sites [79].

Deactivation of arsenic-resistant catalysts occurs at a slower rate, but these catalysts are not 

arsenic proof [79]. Arsenic-resistant catalyst will maintain adequate activity for approximately 

14,000 hours of operation versus approximately 5,500 hours for a nonresistant catalyst [79].

The addition of small amounts of calcium to the fuel has been shown to be an effective 

method of controlling arsenic poisoning [80]. Limestone (CaCO3) can be injected into the flue 

gas to generate the solid Ca3 (AsO4)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 [79]. It should also be noted that 

CaO can be a catalyst poison via formation of CaSO4 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 oxidized to mercuric chloride [HgCl2]). The oxidized 

mercury is water soluble and easier to remove in downstream wet FGD units or SO2 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 stream. The type of coal burned and its associated chlorine 

content is another factor [81].

Studies on simulated flue gas streams and slip streams from actual electric generating

units for multiple catalyst types have shown that mercury oxidation is in the range of 80–90% for 

fresh catalyst and space velocities of approximately 1,000 hr-1, but that the oxidation rate 

declines to 30–80% 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 rate 
Chapter 2 – Selective Catalytic Reduction

2-48

when the space velocity increases. These 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 consistent with the fact that mercury oxidation to HgCl2 is greater at lower temperatures [81].

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 fullscale utility boiler SCR systems [81]. In recent studies, where the NH3 to NOx ratio was varied 

between 0 and 0.95, the suppressive nature of ammonia on mercury oxidation has been 

demonstrated [82]. In the reaction process on the catalyst sites, ammonia adsorption takes 

precedence over the site chlorination, therefore the NOx reduction reaction has precedence over 

the mercury oxidation [47]. Because the favored ammonia adsorption on the catalyst sites 

minimizes the chlorinated sites, the NOx reduction reaction may actually inhibit mercury

oxidation [47]. In several studies, it has been observed that with NOx reduction efficiencies up to 

90%, the mercury oxidation readily occurs [47]. However, at NOx reduction efficiencies beyond 

90%, mercury oxidation is greatly reduced [47].

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 HgCl2 across the SCR catalyst [81]. Coal types, such as 

subbituminous coal, that tend to have lower chlorine levels will have lower HCl concentrations 

in the exhaust gas than bituminous coal, which has higher chlorine levels [81]. Significant 

mercury oxidation by SCR catalysts occurs with bituminous coal, and the oxidation rate is less 

certain with other types of coals [81]. In recent studies, HCl gas was injected directly into the 

flue gas and was varied from 0 to 150 ppmv Cl; the data demonstrated increased mercury

oxidation across the SCR as chlorine was increased [82]. Chlorine appears to have the greatest 

effect on mercury oxidation in the range of 0–50 ppm [47]. The chlorine content does not appear 

to affect the NOx reduction performance of the SCR [47].

There is a close correlation between mercury oxidation on the catalyst and the SO2 to SO3

conversion mechanism on the catalyst [47]. The mercury reaction with HCl and oxygen is 

considered to be a diffusion-controlled reaction that takes place in the gas phase [47]. The 

mercury oxidation rate is faster than the diffusion velocity of mercury through the SCR catalyst 

due to its molecular weight, while the reaction rate of SO2 to SO3 conversion is slower than the 

diffusion velocity of SO2 through the catalyst [47]. As mentioned above in Section 2.2.1 (see 

Catalyst heading), by altering the catalyst to favor mercury oxidation based on this reaction 

mechanism, the active sites oxidize mercury instead of converting SO2 to SO3 [47]. Higher 

mercury oxidation rates of greater than 90% can be achieved while maintaining low SO2 to SO3

conversion rates in the range of less than 0.5% [47].

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 [82]. It was 

determined that there is little to no difference in the mercury oxidation capability of the new and 

regenerated catalysts [82]. 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; slightly higher chlorine content, which would increase 
Chapter 2 – Selective Catalytic Reduction

2-49

mercury oxidation; and lower space velocity), the data showed lower mercury oxidation for the 

aged catalyst [82].

Dioxin/Furan Control

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 commonly used metal oxide catalysts used in SCR for reducing 

NOx 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 NOx and dioxin/furan control, since only one 

pollution control device is needed to reduce 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 and Balance of Plant

Retrofit of SCR on an existing unit has higher capital costs than SCR installed on a new 

system. There is a wide range of SCR retrofit costs due to site-specific factors, scope differences, 

and site congestion [10]. Specific factors that impact the retrofit costs include the following: [10] 

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

 Congestion downstream of the air heater (i.e., buildings, conveyors, existing ESPs, FGD 

system, ID fan, or stack); 

 The age/vintage and manufacturer of the boiler; 

 The design margin of the existing ID fan (i.e., the need to upgrade or replace fan 

impellers, replace ID fans, or add booster fans); 

 The capacity, condition, and design margins of the electrical distribution system; 

 The design margins of the existing structural steel support systems; 

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

 The number, nature, and type of existing items that must be relocated to accommodate

the SCR and associated systems.

While not all of these components are specifically applicable to cement applications, each 

of these issues have an equal counterpart for cement kiln retrofits.

The primary balance of plant impacts include new ID fans and fan foundations 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 dry 
Chapter 2 – Selective Catalytic Reduction

2-50

bottom wall- or tangentially-fired boilers [42]. Differential retrofit costs for an SCR in Germany 

is approximately $200 per MMBtu/hr ($20/kW) [42]. 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 equations 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 costs. 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 capital costs. Boiler operations that have 

varying operating load, frequent startups/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 modifications 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 

requirements [53].

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, SO2, and SO3 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 computational fluid dynamics and chemical 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 uniform 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 high-dust configuration, this involves modeling 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 [42]. 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 based on prior experience and 

computational fluid dynamics and chemical kinetic modeling [33]. For some industrial 

applications, such as cement kilns where flue gas composition varies with the raw materials 
Chapter 2 – Selective Catalytic Reduction

2-51

used, a slip stream pilot study can be conducted to determine whether trace elements and dust 

characteristics of the flue gas are compatible with the selected catalyst. 

This section presents an approach to estimating design parameters that are elements in the 

costing equations used in EPA’s Integrated Planning 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 boilers, it provides sufficient accuracy and detail to be of aid in developing 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, QB, expressed as

MMBtu/hr. It is obtained from the higher heating value, HHV, of the fuel in Btu per pound 

(Btu/lb) and the maximum fuel consumption rate in pounds per hour (lb/hr),

mfuel



:

6

10

1

QB HHV  m

fuel 

(2.3)

Where:

QB, = maximum heat rate input to the boiler, MMBtu/hr

HHV = higher heating value of the fuel, Btu/lb

ṁfuel = maximum fuel consumption rate of the boiler, lb/hr

1/106 = conversion factor of 1 MMBtu/106 Btu

Table 2.5 provides the HHV for various coals that may be used if the HHV for the specific coal 

used in the project is not available.21

Table 2.5: Higher 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 (MMBtu/MWh):

QB BMW  NPHR

(2.4)

 

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

2-52

Where:

BMW = boiler MW rating at full load capacity. 

NPHR = net plant heat rate, MMBtu/MWh

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/MWh) 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 

MMBtu/MWh) can be used [83].

22 Using this value, the heat input rate, QB, for a coal-fired unit

is:

 10 QB BMW

(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 MMBtu/MWh, compared to a typical heat rate of 10 MMBtu/MWh. The 

developers of the costing methodology presented in section 2.4.1 determined that using this 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 equations 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 MMBtu/MWh.

10

NPHR HRF 

(2.6)

Where:

HRF = Heat rate factor

NPHR = net plant heat rate of the system to be costed, MMBtu/MWh

10 = the NPHR that is the basis of the SCR base cost module capital cost, 

MMBtu/MWh.

2.3.3 System Capacity Factor

The total system capacity factor, CFtotal, is a measure of the average annual use of the 

boiler in conjunction with the SCR system. CFtotal is given by:

 

22 In recent years (2003 to 2011), the average NPHR for coal has increased slightly (likely due to aging of 

equipment), and 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

CFtotal = CFplant CFSCR (2.7)

Where:

CFtotal = total system capacity factor

CFplant = boiler capacity, which is the ratio of the actual quantity of fuel burned annually 

to the potential maximum quantity of fuel burned annually

CFSCR = 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, CFplant, is the ratio of actual electric 

output annually to the potential maximum electric annual output, as shown in Equation 2.8a:

8760)



MW

output

plant (B

B

CF

(2.8a)

Where:

BMW = boiler MW rating at full load capacity, MWh. 

Boutput = annual actual MW output, MW/year.

Alternatively, for industrial and utility boilers, the capacity factor of the boiler, CFplant, is 

the ratio of actual quantity of fuel burned annually to the potential maximum quantity of fuel burned 

annually in pounds, i.e.,

mfuel



in lb/hr × 8,760 hr/yr. CFplant is given by:

fuel

fuel

plant maximum annual m

actual annual m

CF







(2.8)

Where:

actual ṁfuel = annual actual fuel consumption rate of the boiler, lb

maximum ṁfuel = 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 factor for the SCR system, CFSCR, is the ratio 

of the actual number of SCR operating days, tSCR, to the total number of plant operating days per 

year:

plant

SCR

SCR t

t

CF 

(2.9)

Where:

tSCR, = actual days of SCR operation annually, days
Chapter 2 – Selective Catalytic Reduction

2-54

tplant = actual days of plant (or boiler) operation in a year, days.

2.3.4 Inlet NOx and Stack NOx

Inlet NOx, represented as NOXin, is the NOx emission level in the flue gas exit stream 

from a boiler prior to the SCR system. Note that NOXin also accounts for combustion controls if 

the boiler is equipped with such controls. The inlet NOx emissions level, obtained from 

analyzing the boiler flue gas stream, is generally given in lb/MMBtu of NO2 [33].

The stack NOx, represented as NOXout is the required NOx emission limit at the stack 

outlet. It is generally set by the plant or regulatory limits and also given in lb/MMBtu of NO2

[33].

2.3.5 NOx Removal Efficiency

The NOx removal efficiency, represented as ηNOx, is determined from the inlet or 

uncontrolled NOx level of the boiler at maximum heat input rate, CFplant =1.0, and the required 

stack emission limit. The equation for the NOx removal efficiency is given by:

x in

x in x out

NO x NO

NO  NO

 

(2.10)

Where:

ηNOx = NOx removal efficiency, fraction

NOxin = inlet NOx level from the boiler, i.e., inlet NOx rate to the SCR, lb/MMBtu (at 

maximum heat input rate, CFplant = 1.0)

NOxout = outlet NOx rate from the SCR, lb/MMBtu

The required NOx removal efficiency is one of the most influential parameters on the overall 

SCR system cost [53]. 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 greater 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 operating) 

average should be obtainable by a power plant boiler with an SCR system.

2.3.6 NOx Removal Rates

The tons of NOx removed annually (ton/yr) are:

NOx Removed/yr = NOxin NOx QB top/ 2,000 (2.11)

Where:

NOx Removed/yr = annual mass of NOx removed by the SCR, tons/yr

QB = maximum heat input rate to the boiler, MMBtu/hr
Chapter 2 – Selective Catalytic Reduction

2-55

top = operating time per year (CFtotal x 8760), hr/yr

2000 = conversion factor for lb/ton.

The pounds of NOx removed per hour (lb/hr) are:

NOx Removed/hr = NOxin NOx QB (2.12)

Where: 

NOx Removed/hr = hourly mass of NOx removed by the SCR, lb/hr

NOxin = inlet NOx of the boiler, lb/MMBtu (at maximum heat input rate, 

CFplant = 1.0)

NOx = NOx removal efficiency of the SCR, expressed as a fraction

QB = 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 NOx 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 NOx 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:

moles of NO moved

moles of reagent injected SRF

x Re



(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 NO2 in the boiler flue 

gas, which requires two moles of NH3 per mole of NO2 instead of one mole of NH3 per mole of 

NO as shown in Equation 2.1a. 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-load flue gas flow rate, including the typical 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, qfluegas, can be calculated using Equation 2.14. 

SCR

fuel B

fluegas

n

q Q T

q

(460 700)

(460 )







(2.14)
Chapter 2 – Selective Catalytic Reduction

2-56

Where:

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

(acfm)

qfuel = base case flue gas volumetric flow rate factor, ft3/min-MMBtu/hr

T = operating gas temperature at the inlet to the SCR, °F

nSCR = 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.

“Base 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 using procedures in Reference [84] for typical coals, 

typical boiler excess air levels (i.e., 20 %), and typical SCR flue gas conditions (–10 inches 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, Appendix A-7.

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

Coal Type

Estimated value of qfuel

(ft3/min-MMBtu/hr)

Eastern Bituminous 484

Powder River Basin 516

Lignite 547

Note that in general, the number of reactors, nSCR, 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 developed 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 

costs are essentially identical.

2.3.9 Space Velocity and Area Velocity

The space velocity, Vspace, is defined as the inverse of the residence time and is given by:

ResidenceTime

Vspace

1



(2.15)

Where:

Vspace = the volumetric 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, hr
Chapter 2 – Selective Catalytic Reduction

2-57

Space velocity is calculated from the experimentally measured flue gas volumetric flow rate at 

the reactor inlet, represented as qfluegas, and the reactor/catalyst volume, represented as Volcatalyst, 

given by:

catalyst

fluegas

space Vol

q

V 

(2.16)

Where:

Volcatalyst = volume of the reactor or catalyst layers, ft3

SCR system designers and vendors use the concept of area velocity, Varea, 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, Aspecific, as follows:

specific

space

area A

V

V 

(2.17)

Where:

Varea = the space velocity divided by the catalyst pore surface area, ft3/ft2·hr

Aspecific = the specific surface area of the catalyst divided by the catalyst volume, ft2/ft3

Aspecific is given in units of length2/length3, is sometimes referred to as the contact surface area of 

the catalyst, and must be provided by the catalyst manufacturer. 

2.3.10 Theoretical NOx Removal Efficiency

Equation 2.10 defines the NOx removal efficiency. However, in SCR, NOx removal efficiency 

changes with catalyst activation. The following theoretical equation allows for estimation of

removal efficiency, ηNOx, based on the catalyst activity constant, Kcatalyst, at a specified

time, t [33]. The theoretical NOx removal efficiency is:

NOx = SFR (1-e

a

) (2.18)

where















 



space

catalyst specific

V

K A

a

(2.19)

Where:

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

Both Kcatalyst and Aspecific are typically provided by the catalyst manufacturer.
Chapter 2 – Selective Catalytic Reduction

2-58

According to this equation, the NOx removal efficiency increases with increasing 

NH3/NOx 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 NOx 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. Equations 2.16, 2.18, and 2.19 can be combined 

and rearranged to determine the theoretical catalyst volume [56]. Substituting the definition of 

space velocity (Equation 2.16) into the definition of a (Equation 2.19), and then substituting that 

into Equation 2.18 and solving for the volume of the catalyst gives:

catalyst specific

NO x

fluegas

catalyst K A

SRF

q

Vol















































  





ln 1

(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:

SCR

adj

catalyst B adj adj x adj adj N

T

Vol  2.81 Q   Slip  NO  S 

(2.22)

where NSCR isthe number of SCR reactors and the adjustment factors include:

 NOx efficiency adjustment factor, adj:

0.2869 (1.058 ) adj    NOx

(2.23)

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

Slip 1.2835 (0.0567 Slip) adj   

(2.24)

 NOx adjustment factor for inlet NOx, NOxadj:

  NOxadj NOxin

 0.8524  0.3208

(2.25)

 Sulfur in coal adjustment factor, Sadj:

S 0.9636 (0.0455 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, Tadj, for gas temperatures other than 700°F (370°C):

   

5 2

Tadj 15.16  0.03937T  2.7410 T



(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 equation for the catalyst cross-sectional area, 

Acatalyst, is given by:



























min

60 sec

sec

16 ft

q

A

fluegas

catalyst

(2.28)

Where:

Acatalyst = cross-sectional area of the catalyst, ft2

The SCR reactor cross-sectional area, ASCR, is approximately 15% greater than the 

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

ASCR = 1.15 x Acatalyst (2.29)

Where:

ASCR = cross-sectional area of the SCR reactor, ft2

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 

screening 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:

 

1/ 2

w ASCR l  

(2.30)

Where:

w = width of the SCR reactor, ft

l = length of the SCR reactor, ft

An initial value for the number of catalyst layers is estimated first. This estimate is then 

checked by calculating the catalyst height for each layer. The initial estimate for the number of 
Chapter 2 – Selective Catalytic Reduction

2-60

catalyst layers can be determined from the total 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, nlayer, is:

layer catalyst

catalyst

layer h' A

Vol

n





(2.31)

Where:

nlayer = number of catalyst layers

h′layer = nominal height of each catalyst layer, ft

This value of nlayer 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, hlayer, 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:

1

n A

Vol

h

layer catalyst

catalyst

layer 





















(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, ntotal, includes all empty catalyst layers that will be 

installed:

ntotal = nlayer + nempty (2.33)

Where:

ntotal = total number of catalyst layers

nempty = 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 nlayer

in Equation 2.31 and hlayer in Equation 2.32 for a particular system. The optimum range of values for h′layer differs 

depending on the inlet NOx rate. For example, for a relatively high inlet NOx rate of 0.86 lb/MMBtu, the optimum 

range of h′layer is 3.1 feet to 4.2 feet. The optimum range shifts to smaller values when the inlet NOx rate is lower, 

but the range is expected to include 3.1 feet for all inlet NOx rates greater than 0.1 lb/MMBtu. Higher values of 

h′layer 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 hlayer. 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, hSCR, including the initial and future catalyst layers, the flowrectifying layer, space for soot blowers and catalyst loading, but excluding the inlet and outlet 

ductwork and hoppers, is determined from the equation:

hSCR = ntotal (c1 + hlayer) + c2 (2.34)

Where:

hSCR = height of the SCR reactor, ft

c1 = constant based on common industry practice, i.e., 7 ft, ft

c2 = constant based on common industry practice, i.e., 9 ft, ft

where the constants are based on common industry practice of c1 = 7 ft and c2 = 9 ft.

2.3.13 Estimating Reagent Consumption and Tank Size

The rate of reagent consumption or mass flow rate of the reagent, 

m



reagent, generally 

expressed as pounds per hour (lb/hr), can be calculated using the inlet NOx in lb/MMBtu and 

heat input rate, QB, in MMBtu/hr.

x

in

NO

x B NOx reagent

reagent M

NO Q SRF M

m

   





 (2.35)

Where:

ṁreagent = mass flow rate, or consumption rate, of the reagent, lb/hr

Mreagent = the molecular weight of the reagent (60.06 pounds per mole [lb/mole] for urea, 

17.03 lb/mole for ammonia)

MNOx = the molecular weight of NO2 (46.01 lb/mole).

NOx = NOx removal efficiency of the SCR, expressed as a fraction

The molecular weight of NO2 is used because the NOx emissions, NOxin, are given in lb/MMBtu 

of NO2.

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

msol

 , is given by:

sol

reagent

sol C

m

m



 

(2.36)

Where:

ṁsol = mass flow rate of the aqueous reagent solution, lb/hr

Csol = the concentration of the aqueous reagent solution, by weight fraction.

The solution volume flow rate, qsol, generally expressed as gallons per hour (gph), is:
Chapter 2 – Selective Catalytic Reduction

2-62   7.4805

sol

sol

sol

m

q





(2.37)

Where:

qsol = solution volume flow rate, gph

 sol = the density of the aqueous reagent solution, lb/ft3

7.4805 = conversion factor of 7.4805 gal/1 ft3.

The ρsol is 56.0 lb/ft3for a 29% solution ammonia and 71.0 lb/ft3for 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, tstorage, is:

Voltank = qsol × tstorage × 24 (2.38)

Where:

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

tstorage = 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 onsite storage requirement is for 14 days of SCR 

operation.

2.4 Cost Analysis 

The cost-estimating methodology presented 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 2012 dollars (2012$).

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 installation cost data, the IPM equations 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.chemengonline.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 difference between the IPM methodology and the 

methodology presented here is that the IPM methodology includes owner’s costs (for owner 

activities related to engineering, 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), owner’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 applicable 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 capacity 

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. The TCI equations 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. The equations may overestimate costs for some simple retrofits of existing 

plants. For retrofits that are more complicated than average, a retrofit factor of greater than 1 can 

be used to estimate capital costs 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 The catalyst is a ceramic honeycomb with an 

operating life of 3 years at full load operations [33]. The cost equations are sufficient for NOXout

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 [52]. High-dust units tend to have higher capital costs

when there are space constraints while tail-end units tend to have higher operating costs [52]. 

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 elements 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 estimating capital costs is based on an engineering and design firm’s inhouse 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 minority 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 procedures 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. Costs 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, land, 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 boilers 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 for 

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 for 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 unit 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:

PELEV

P

ELEVF 0

(2.39a)

Where:

ELEVF = elevation factor

P0 = atmospheric pressure at sea level, 14.7 pounds per square inch absolute (psia)

PELEV = 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, PELEV can be calculated using equation 2.39b[108]:

144

1

518.6

59 (0.00356 ) 459.7

2116

5.256

 









   

 

h

PELEV

(2.39b)

Where:

PELEV = atmospheric pressure at elevation of the unit, psia

h = altitude, feet

2.4.1.1 Utility Boilers (Coal-fired)

Utility, coal-fired units. The capital cost equation for coal-fired units (applicable for ≥25 

MW) is:

TCI SCR RPC APHC BPC 1.3 Cost   

(2.40)

Where:

TCI = total capital investment for a SCR on a coal-fired boiler, $

SCRCost = cost of the SCR, $
Chapter 2 – Selective Catalytic Reduction

2-66

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 complicated 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 costs 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:

SCRCost NRF BMW  HRF CoalF  ELEVF  RF 0.2 0.9 2 270,000

(2.41)

Where:

SCRCost = SCR unit costs, $

270,000 = constant in the equation

NRF = NOx removal factor (ηNOx/80)

BMW = electric generating capacity of the unit supplied by the boiler, MW

HRF = heat rate factor

CoalF = coal factor (CoalF=1 if bituminous; CoalF=1.05 if PRB; CoalF=1.07 if Lignite)

ELEVF = elevation factor.

RF = retrofit factor (RF = 0.8 for new construction; RF = 1 for retrofits of average 

difficulty)

The NOx 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 NOx removal efficiency (ηNOx) 

divided by 80. The CoalF is 1 for bituminous coal, is 1.05 for Powder River Basin (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 NOx 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 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:

RPC  NOxin  BMW  NPHR NOx   RF 0.2 5 490,000 

(2.42)

Where:

RPC = Reagent preparation cost, $

490,000 = constant in the equation

NOxin = inlet NOx level from the boiler, i.e., inlet NOx rate to the SCR,

lb/MMBtu
Chapter 2 – Selective Catalytic Reduction

2-67

NPRH = net plant heat rate, MMBtu/MWh

ηNOx = NOx 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 SO3 control is necessary. An air pre-heater 

modification is necessary for the control of SO3 for boilers that burn bituminous coal where the 

SO2 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 SO3 formation. If lower sulfur content coal types are used, then no air pre-heater 

modification is needed. The air pre-heater modification costs are calculated as follows:

APHC  BMW  HRF CoalF  AHF  RF 0.7 8 69,000

(2.43)

Where:

APHC = Air pre-heater cost, $

69,000 = constant in the equation

AHF = air heater factor (AHF=1 if bituminous coal and SO2 ≥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 SO2 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:

BPC  BMW  HRF CoalF  ELEVF  RF 0.4 2 460,000

(2.44)

Where:

BPC = Balance of plant cost, $

460,000 = constant in the equation

ELEVF = elevation factor.

RF = retrofit factor (RF = 0.8 for new construction; RF = 1 for retrofits of average 

difficulty)

2.4.1.2Utility Boilers (Oil- and Gas-fired)
Chapter 2 – Selective Catalytic Reduction

2-68

Utility, oil- and gas-fired units (≥25 MW to 500 MW). The capital cost equation for oiland gas-fired units, applicable to ≥25 MW to 500 MW, is:

B ELEVF RF

B

TCI MW

MW

   















 

0.3 5

200 80,000

(2.45)

Where:

TCI = total capital investment for an SCR unit on an oil-fired or gas-fired boiler, $

80,000 = installed cost of an SCR system in 2012$ for a 200 MW oil- or gas-fired boiler, 

$/MW

BMW = electric generating capacity of the unit supplied by the boiler, MW

ELEVF = 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 oiland gas-fired units, unlike the 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, oil- and gas-fired units (>500 MW). For oil- and gas-fired units >500 MW, the 

normalized costs in $/kW are assumed to be equivalent to the costs for a 500 MW boiler unit. 

Thus, the equation for units >500 MW reduces to:

TCI  60,670BMW ELEVF RF

(2.46)

With respect to high-dust and tail-end SCR units, high-dust units typically require larger catalyst 

volume that increases 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 m3/MW, and tail-end units require less than half the catalyst 

volume of a high dust unit (or less than 0.75 m3/MW) [85]. The lower catalyst volume for tailend units helps reduce initial 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 costs 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 costs 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 Industrial Boilers (Coal-fired)

Industrial, coal-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 the costs for industrial boilers since current retrofits of utility 

boilers are likely more complex than for industrial boilers. The capital cost equation is:

TCI SCR RPC APHC BPC 1.3 Cost   

(2.47)

Where:

TCI = total capital investment for a SCR on a coal-fired boiler, $

SCRCost = 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 costs 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:

SCRCost NRF  QB CoalF  ELEVF  RF 0.2 0.9 2 270,000 0.1

(2.48)

Where:

SCRCost = SCR unit costs, $

270,000 = constant in the equation

NRF = NOx removal factor (ηNOx/80)

QB = maximum heat rate input to the boiler, MMBtu/hr 

CoalF = coal factor (CoalF=1 if bituminous; CoalF=1.05 if PRB; CoalF=1.07 if Lignite)

ELEVF = elevation factor.

RF = retrofit factor (RF = 0.8 for new construction; RF = 1 for retrofits of average 

difficulty)

The NOx 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 NOx removal efficiency (ηNOx)

divided by 80. The CoalF is 1 for bituminous coal, is 1.05 for Powder River Basin (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 into the SCR are based on the NOx 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 
Chapter 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:

RPC  NOxin QB

nNOx   RF 0.2 5 490,000

(2.49)

Where:

RPC = Reagent preparation cost, $

490,000 = constant in the equation

NOxin = inlet NOx level from the boiler, i.e., inlet NOx rate to the SCR, lb/MMBtu

ηNOx = NOx 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 preheater modification costs are included only where SO3 control is necessary. An air pre-heater 

modification is necessary for the control of SO3 for boilers that burn bituminous coal where the 

SO2 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:

APHC   QB CoalF  AHF  RF 0.7 8 69,000 0.1

(2.50)

Where:

APHC = Air pre-heater cost, $

69,000 = constant in the equation

CoalF = coal factor (CoalF=1 if bituminous; CoalF=1.05 if PRB; CoalF=1.07 if Lignite)

AHF = air heater factor (AHF=1 if bituminous coal and SO2 ≥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 SO2 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:

BPC   QB CoalF  ELEVF  RF 0.4 2 460,000 0.1

(2.51)
Chapter 2 – Selective Catalytic Reduction

2-71

Where:

BPC = Balance of plant cost, $

460,000 = constant in the equation

ELEVF = elevation factor.

RF = retrofit factor (RF= 0.8 for new construction; RF= 1 for retrofits of average 

difficulty)

2.4.1.4 Industrial Boilers (Oil- and Gas-fired)

Industrial, oil-fired units (≥275 to ≤5,500 MMBtu/hr). The capital cost equation for oilfired industrial boilers was developed by modifying equation 2.45. The 200 MW rating of the 

generating unit for the base utility boiler in equation 2.45 was converted to a heat input capacity 

of 2,200 MMBtu/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 2.45 by the NPHR of 11,000 

Btu/kwh gives an estimated cost of 7,270 $/MMBtu/hr. The resulting equation, applicable to oilfired industrial boilers with a heat input capacity of ≥275 to ≤5,500 MMBtu/hr, is:

Q ELEVF RF

Q

TCI B

B

   















 

0.3 5

2,200 7,270

(2.52)

Where:

TCI = total capital investment for SCR on an oil-fired boiler, $

7,270 = estimated installed cost of an SCR unit in 2012$ for an oil-fired boiler that has a 

2,200 MMBtu/hr rating at full load capacity, $/MMBtu/hr

QB = maximum heat rate input to the boiler, MMBtu/hr

ELEVF = 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, gas-fired units (≥205 to ≤4,100 MMBtu/hr). The capital cost equation for gasfired industrial boilers was developed 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 ≥205 to ≤4,100 MMBtu/hr, is:
Chapter 2 – Selective Catalytic Reduction

2-72 Q ELEVF RF

Q

TCI B

B

   















 

0.3 5

1,640 9,760

(2.53)

Where:

TCI = total capital investment for SCR on a gas-fired boiler, $

9,760 = estimated installed cost of an SCR unit in 2012$ for a gas-fired boiler that has a 

1,640 MMBtu/hr rating at full load capacity, $/MMBtu/hr

QB = maximum heat rate input to the boiler, MMBtu/hr

ELEVF = 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 gas-fired units may 

overestimate the costs.

Industrial, oil-fired units (>5,500 MMBtu/hr). For oil-fired industrial boilers >5,500 

MMBtu/hr, the normalized costs in $/MMBtu/hr are assumed to be equivalent to the costs for a 

5,500 MMBtu/hr boiler unit. Thus, the equation for units >5,500 MMBtu/hr reduces to:

TCI  5,275QB ELEVF  RF

(2.54)

As noted earlier, applying the ELEVF to the TCI equation for industrial oil-fired units may 

overestimate the costs.

Industrial, gas-fired units (>4,100 MMBtu/hr). For gas-fired industrial boilers >4,100 

MMBtu/hr, the normalized costs in $/MMBtu/hr are assumed to be equivalent to the costs for a 

4,100 MMBtu/hr boiler unit. Thus, the equation for units >4,100 MMBtu/hr reduces to:

TCI  7,082QB ELEVF  RF

(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 by 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 [86]. Each of these costs is 

discussed in the sections below. A more detailed discussion 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 annual heat 

input rate of the boiler and SCR system using CFtotal.. 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 individually below. Equations for these variable cost items were 

derived in Reference [8].

Operating costs also result from small decreases in boiler efficiency due to operation 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 operating cost is 

incurred for the steam or electric power used for compressed air, as required for the relatively 

infrequent operation of soot blowers. These operating costs are generally small and site-specific. 

Therefore, they are not discussed in this report.

























































































Cost

Catalyst

Annual

Cost

Electricit y

Annual

Cost

Reagent

Annual

Cost

Maintenance

Annual

DAC

(2.56)

Operating and Supervisory Labor

In general, operation of an SCR system requires only minimal, operating or supervisory 

labor. The SCR reactor is a stationary device with no moving parts. Further, the SCR system 

incorporates only a few pieces of rotating equipment (e.g., pumps, motors). The IPM [8] 

estimates operating labor time as 4 hours per day.

Maintenance

The annual maintenance labor and material cost in dollars per year ($/yr), including 

nozzle tip replacement for the injectors, is assumed to be 0.5% of the TCI in dollars [7a, 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, qsol, the operating time per year, top, and the cost of reagent in dollars per gallon, Costreag:

Annual Reagent Cost = qsol x Costreag x top (2.58)

where qsol is in gallons per hour (gph). The example here is for use of an aqueous ammonia 

solution, i.e., qsol, however, if the cost of the ammonia is available in $/lb of ammonia, then the 

equation to estimate reagent costs would be based on mreagent from Equation 2.35 instead of qsol.

The operating time per year, top, is estimated using the capacity factor, CFtotal:

t

op

 CFtotal 8760

(2.59)

Utilities

Power consumption for utility boilers. The electrical power consumption, P, in kW 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] :

0.4 3 P B (1,000) (0.0056) (CoalF HRF) 

MW    

(2.60)

Where:

P = electrical power consumption of the SCR system, kW

1,000 = conversion factor for 1,000 kW/MW

0.0056 = 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 kW is 

estimated for 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:

0.4 3 P (0.1 Q ) (1,000) (0.0056) (CoalF HRF)   B    

(2.61)

The annual cost of electricity is estimated from the equation:

Annual Electricity Cost = P x Costelect x top (2.62)

Where:

Costelect = 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 highdust and low-dust SCR is shown in Table 2.8. As an example, one source cites a total power 
Chapter 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 [85]. 

Depending on the site configuration and space constraints, application of high-dust SCR 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 [85].

Table 2.8: Comparison of Power Consumption for High-Dust and Low-Dust SCR [85]

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 increases 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 typically 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 annually, and the future worth of the 

catalyst. This cost methodology assumes a guaranteed 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 methodology employed for power plants in the IPM [8].

Under catalyst replacement cost methodology 1, if 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 been 
Chapter 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 = nSCR Volcatalyst

CCreplace

Rlayer

(2.63)

Where:

Catalyst Replacement

Cost = cost to replace the SCR catalyst, either replacing all catalyst or 

replacing 1 layer at a time, $

Volcatalyst = volume of the reactor or catalyst layers, ft3

nSCR = number of SCR reactor chambers

CCreplace = cost of catalyst, dollars per cubic foot ($/ft3)

Rlayer = catalyst replacement factor (Rlayer = 1 for full replacement and Rlayer = 

nlayer for replacing 1 layer at a time)

The catalyst cost, CCreplace, should reflect the current costs for the catalyst, 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 be 

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:

FWF

Cost

Replacement

Catalyst

Annual Catalyst Replacement Cost 





















 (2.64)

where FWF is the future worth factor. Future worth is used because the annual catalyst 

replacement cost is accrued starting in the first year of operation, while catalyst replacement 

purchases occur every few years. To account for the time value of money, the FWF amortizes the 

catalyst cost 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:

 













 



1 1

1

Y

i

FWF i

(2.65)

Where:

i = interest rate, fraction
Chapter 2 – Selective Catalytic Reduction

2-77

Y = term, years.

The term, Y, is given by the equation:

year

catalyst

h

h

Y 

(2.66)

Where:

hcatalyst = operating life of the catalyst, hours

hyear = 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, the 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]:

( ) (0.4) ( ) ( ) ( ) 35.3

Re

2.9 0.7 1

      















BMW CoalF NRF CCreplace placement Cost

Annual Catalyst

(2.67)

Where:

Annual Catalyst Replacement Cost = cost to replace the SCR catalyst, $/yr

CCreplace = cost of catalyst, dollars per cubic meter ($/ft3)

35.3 = conversion factor for $/ft3to $/m3.

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 poisons 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 [52]. In addition, 

concentrations of SO2 in the flue gas are low following the wet scrubber and there are fewer 

concerns related to SO3 formation and ammonium salt deposition [52].

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 cites 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 [89], and another source reports tail-end catalysts that lasted for 

100,000 operating hours [52].

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 based 
Chapter 2 – Selective Catalytic Reduction

2-78

on the anticipated equipment lifetime28 and the annual interest rate employed.29 For the purposes 

of this cost example, the equipment lifetime of an SCR system is assumed to be 30 years for

power plants and 20 to 25 years for industrial boilers. These assumptions 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 lawsuit against TVA coalfired electric generation units indicated expected 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 [92]. 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 can 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 [41]. The cost of 

overhead for an SCR system is also considered to be zero. An SCR system is not viewed as riskincreasing 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 [41].

Finally, there are two categories of overhead, payroll and plant. Payroll overhead includes 

expenses related to labor employed in operation 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, consist of both 

administrative charges and capital recovery, which can be expressed as:



































ery

Capital

Charges

Administrative

Indirect Annual Cost

Recov

(2.68)

Administrative Charges

Administrative charges may be calculated as:

Administrative Charges = 0.03 × ((

Operator

Labor Cost) + 0.4 × (

Annual Maintenance

Cost )) (2.69)

Where 

 

28 The term “equipment life” as used here in this chapter and through the 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. Available 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/Proposed-Rules/regxx/noxreclaim_wgm_042915b.pdf?sfvrsn=2.
Chapter 2 – Selective Catalytic Reduction

2-79

Operator Labor Cost = tSCR 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 requires only minimal, operating or supervisory labor.

Capital Recovery

Capital recovery is estimated as:

CR = CRF × TCI (2.70)

where TCI is the total investment, and CRF is the capital recovery factor and defined by:

 

1  1

1

 





n

n

i

i i

CRF

(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:













































Cost

Annual

Indirect

Cost

Annual

Direct

Total Annual Cost

(2.72)

Cost Effectiveness

The cost in dollars per ton of NOx removed per year is:

Removed/yr NOx

TAC Cost Effectiveness 

(2.73)

Where:

Cost Effectiveness = the cost effectiveness, $/ton

NOx Removed/yr = annual mass of NOx removed in the SCR, ton/yr

2.5 Example Problem #1 – Utility Boiler

An example problem that calculates both the design parameters and capital and annual 

costs for an SCR system applied to a 120 MW utility boiler firing bituminous coal is presented 

below. The following assumptions are made to perform the calculations:

Fuel High Heating Value, HHV 12,000 Btu/lb

Net Plant Heat Rate, NPHR 10 MMBtu/MWh

Maximum Actual Output 102 MW
Chapter 2 – Selective Catalytic Reduction

2-80

Total Annual Fuel Consumption, actual mfuel 7.45 x 108lb

Number of plant (boiler) operating days, tPlant 365 days

Number of SCR operating days, tSCR 365 days

Inlet NO Level, NOXin 0.35 lb/MMBtu

Required Annual Average Controlled NOx Emission Level, NOXout 0.05 lb/MMBtu

Acceptable Ammonia Slip, Slip 2.0 ppm

Base Case Flue Gas Flow Rate Factor, Eastern Bituminous, qfuel 484 ft3/min per MMBtu/hr

Fuel Sulfur Content, S 1.0% by weight

Stoichiometric Ratio Factor for Ammonia, SFR 1.05

Stored Ammonia Concentration, Csol 29%

Number of Days of Storage for Ammonia, t 14 days

Pressure Drop for SCR Ductwork, ∆Pduct 3 inches w.g.

Pressure Drop for each Catalyst Layer, ∆Pcatalyst 1 inch w.g.

Temperature at SCR Inlet, T 650°F

Plant elevation, PELEV 1,500 ft

In addition to these assumptions, the estimated economic factors for the cost equations are:

Cost year 2012$

Equipment Life 30 years

Annual Interest Rate 7%

Catalyst Cost31 5,500 $/m3($160/ft3) [7b] 

Electrical Power Cost31 0.0372 $/kWh [93] 

29% Ammonia Solution Cost31 0.475 $/lb of ammonia [average of vendor quotes 

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 #132

Boiler Calculations

The boiler annual heat input rate, QB, is calculated 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,

mfuel



:

 

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 operating 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 utility 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 imply 

greater accuracy of the numbers.
Chapter 2 – Selective Catalytic Reduction

2-81

hr

MMBtu

MMBtu

Btu

hr

lb

lb

Btu

QB1,200

10

12,000 100,000

6







The plant capacity factor is calculated from the maximum and annual MW output using 

Equation 2.8a:

0.85 85%

120

102

  

MW

MW CFplant

The SCR system capacity factor is calculated from the months of SCR operation, 

12 months, using Equation 2.9:

1.0 100%

365

365

  

days

days CFSCR

The total capacity factor including both plant and SCR capacity factors is given by:

CFtotal = 0.85 x 1.0 = 0.85 = 85%

The flue gas flow rate using Equation 2.14 is:

acfm

F

F

hr

MMBtu

hr

MMBtu min

ft

q flue gas 556,000

(460 700 ) 1

1,200 (460 650 )

( )

484 3



 

  

 



The NOx removal efficiency, ηNOx, is calculated from the inlet NOx level and the required

controlled NOx emission level using Equation 2.10:

0.857 85.7%

0.35

0.35 0.05

 





MMBtu

lb

MMBtu

lb

MMBtu

lb

 N Ox

SCR Reactor Calculations

The catalyst volume using Equation 2.22 and the equations for each adjustment factor is:

Volcatalyst = 2.81 x 1,200 MMBtu/hr

× [0.2869 + (1.058 x 0.86)] (adjx)

× [0.8524 + (0.3208 x 0.35)] (NOxadj)

× [1.2835 – (0.0567 x 2.0)] (Slipadj)
Chapter 2 – Selective Catalytic Reduction

2-82

× [0.9636 + (0.0455 x 1.0)] (Sulfuradj)

× [15.16 – (0.03937 x 650) + (0.0000274 x 6502)] (Temperatureadj)

= 5,300 ft3

The catalyst and SCR cross-sectional areas using Equations 2.28 and 2.29 are:

2

579

min

60 16

555,766 ft

s

s

ft

acfm Acatalyst 



























2 2

A 1.15 579 ft 666 ft SCR  

The length and width of the reactor using Equation 2.30 is:

l w 666 25.8 ft 1 2  

The first estimate of the number of catalyst layers using Equation 2.31 is:

2.9

3.1 579

5,300





nlayer 

Rounding this value gives, nlayer = 3.

Checking the actual catalyst height using Equation 2.32:

1 4.0

3 579

5,300

2

3

 





ft

ft hlayer

This value is within the design height limits of 2.5–5 feet.

The total number of catalyst layers is determined by Equation 2.33 with one empty 

catalyst layer:

ntotal = 3 + 1 = 4

The SCR height, excluding the outlet duct and hoppers using Equation 2.34 is:

hSCR = 4 × (7 + 4.0) + 9 = 53 ft
Chapter 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 NO2, 46.01g/mole. For an SRF of 1.05, the reagent mass flow rate is given by 

Equation 2.35:

hr

lb

mole

lb

mole

lb

hr

MMBtu

MMBtu

lb

m

N Ox

reagent 140

46.01

0.35 1,200 0.857 1.05 17.03



   

 

The mass flow rate of 29% aqueous ammonia solution is given by Equation 2.36:

hr

hr lb

lb

msol 482

0.29

140

  

The solution volume flow rate can then be calculated from Equation 2.37 where sol is the

density of the 29% aqueous ammonia solution, 56.0 lb/ft3at 60oF, and the conversion factor is 

7.481 gal/ft3:

gph

ft

lb

ft

gal

hr

lb

qsol 64

56.0

482 7.481

3

3







The 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 requirement is given by Equation 2.38:

  gal

day

hr Vol gph days tank 21,600 24 64 14 















  

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):

1.06

13.9

14.7

 

psia

psia ELEVF
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 SCRcost, RPC, APHC and BPC must 

be calculated individually using equations 2.41, 2.42, 2.43, and 2.44, respectively. These 

calculations are shown below.

1.00 1.06 1 $23,652,228

10

10 /

120

80

85.7%

270,000

0.2 0.9 2

   











   











 

MMBtu MWh SCRCost MW 490,000 0.35 120 10 0.857 1 $2,134,380

0.2 5

  



























   

MWh

MMBtu MW

MMBtu

lb RPC 69,000 120 1 1 0 1 $0

0.7 8 APHC   MW     

460,000 120 1 1 1.06 1.4 $3,628,130 0.4 2 BPC   MW     

TCI 1.3$23,652,228$2,134,380$0$3,628,130  $38,239,160

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 2.57:

Annual MaintenanceCost  0.005$38,239,160  $191,196/ yr

An estimate for power consumption is given by Equation 2.60:

P (120 MW) (1,000 k W / MW) (0.56/100) (1 1) 670 k W 0.4 3     

Electricity cost can then be estimated from Equation 2.62:

yr kWh yr

hr Annual Electricit yCost k W $0.0372 $185,584 0.85 8,760

 670    

Reagent cost is estimated using Equation 2.58:
Chapter 2 – Selective Catalytic Reduction

2-85

yr lb yr

hr

hr

lb Annual Reagent Cost $ $499,237

140 8,760 0.8570.475 

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 

methodology 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, Rlayer = 3, and the total 

catalyst replacement cost is given by:

$273,232

3

$160 5,300 3

3







ft

ft

Catalyst Replacement Cost

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 hyear is 8,760 hr/yr, and Y is given by:

2.7 3

8,760

24,000

  

hours

hours Y

And the future worth factor for the catalyst replacement is given by:

 

0.3111

1 0.07 1

1

0.07 3

 

FWF  

The annual catalyst replacement cost can then be estimated using Equation 2.64:

 

yr

Annual Catalyst placement Cost $84,989 Re  0.3111 $273,232 

Alternatively, using catalyst replacement cost methodology 2, the annual catalyst 

replacement cost can be estimated using Equation 2.67:

m yr

ft

ft

MW

placement Cost

Annual Catalyst $160 35.3 $277,254

80

85.7

(120 ) (0.4) (1)

Re 3

3

3

0.7 1

2.9

































 











    















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

yr yr yr yr yr

Total Direct AnnualCost $191,196 $185,584 $499,237 $84,989 $961,000    

Alternatively, if using cost methodology 2 for the catalyst replacement cost, the total 

direct annual cost is:

yr yr yr yr yr

Total Direct AnnualCost $191,196 $185,584 $499,237 $277,254 $1,153,300    

As discussed in section 2.4.2, property taxes and overhead are both assumed to be zero, 

and insurance costs are assumed to 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 2.69 as:

  yr

day day

hours

year

days AdministrativeCharges 0.4 $191,196 $4,922/365 4 $60 0.03 

















  















   

The capital recovery factor, CRF, is defined by Equation 2.71 as:

 

 

0.0806

1 0.07 1

0.07 1 0.07

3 0

3 0



 



CRF 

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 2.68:

yr yr yr

IDAC $4,922 $3,081,556 $3,086,478  

The total annual cost is the sum of the direct annual and indirect annual costs given by 

Equation 2.72. If using cost methodology 1 for the catalyst replacement cost, the total annual 

cost is:

yr yr yr

Total Annual Cost $961,000 $3,086,478 $4,047,478  

Alternatively, if using cost methodology 2 for the catalyst replacement cost, the total 

annual cost is:

yr yr yr

Total Annual Cost $1,153,300 $3,086,478 $4,239,778  
Chapter 2 – Selective Catalytic Reduction

2-87

The annual cost in terms of NOx removed can be calculated using the total annual cost 

and the tons of NOx removed annually. The annual reduction in NOx emissions is given by 

Equation 2.11:

tons yr

ton

lb

yr

hr

hr

MMBtu

MMBtu hr

lb

NO Removed yr x1,340 /

2,000

0.857 1,200 0.85 8,760

/

0.35

/ 

   



and the cost effectiveness is estimated using Equation 2.73. If using cost methodology 1 for the 

catalyst replacement cost, the cost effectiveness is:

tons ton

yr Costof NO Removal x

$3,020

1,340

$4,047,478

 

Alternatively, if cost methodology 2 is used for the catalyst replacement cost, then the 

cost effectiveness is:

tons ton

yr Costof NO Removal x

$3,164

1,340

$4,239,778

 

2.6 Example Problem #2 – Industrial Boiler

An example problem that calculates both the design parameters 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 12,000 Btu/lb

Maximum Fuel Consumption Rate, 

m



fuel 4.58 x 104

lb/hr

Total Annual Fuel Consumption, actual mfuel 3.30 x 108lb

Number of plant boiler operating days 333 days

Number of SCR operating days, tSCR 333 days

Inlet NO Level, NOXin 0.35 lb/MMBtu

Required Annual Average Controlled NOx Emission Level, NOXout 0.05 lb/MMBtu

Acceptable Ammonia Slip, Slip 2.0 ppm

Base Case Flue Gas Flow Rate Factor, Eastern Bituminous, qfuel 484 ft3/min per MMBtu/hr

Fuel Sulfur Content, S 1.0% by weight

Stoichiometric Ratio Factor for Ammonia, SRF 1.05

Stored Ammonia Concentration, Csol 29%

Number of Days of Storage for Ammonia, t 14 days

Pressure Drop for SCR Ductwork, ∆Pduct 3 inches w.g.

Pressure Drop for each Catalyst Layer, ∆Pcatalyst 1 inch w.g.

Temperature at SCR Inlet, T 650°F

Plant elevation, PELEV <500 ft above sea level
Chapter 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 30 years

Annual Interest Rate 7%

Catalyst Cost33 5,500 $/m3($160/ft3) [7b] 

Electrical Power Cost33 0.067 $/kWh [93] 

29% Ammonia Solution Cost33 0.475 $/lb of ammonia [average of vendor quotes 

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 #234

Boiler Calculations

The boiler annual heat input rate, QB, is calculated 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, mfuel



:

hr

MMBtu

MMBtu

Btu

hr

lb

lb

Btu

QB 550

6

10

12,000 45,800







The plant capacity factor is calculated from the maximum and annual average fuel 

consumption:

0.82 82%

8,760

hr

lb 4.58 10

3.30 10

4

8

 



















yr

hr

yr

lb

CFplant

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 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 utility bill, 

while electricity 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 1.0 100%

333

333

  

days

days CFSCR

The total capacity factor including both plant and SCR capacity factors is given by:

CFtotal = 0.82 x 1.0 = 0.82 = 82%

The flue gas flow rate using Equation 2.14 is:

acfm

F

F

hr

MMBtu

hr

MMBtu min

ft

qfluegas 255,000

(460 700 ) 1

550 (460 650 )

( )

484 3



 

  

 



The NOx removal efficiency, ηNOx, is calculated from the inlet NOx level and the required 

controlled NOx emission level using Equation 2.10:

0.857 85.7%

0.35

0.35 0.05

 





MMBtu

lb

MMBtu

lb

MMBtu

lb

 N Ox

SCR Reactor Calculations

The catalyst volume using Equation 2.22 and the equations for each adjustment factor is:

Volcatalyst = 2.81 x 550 MMBtu/hr

× [0.2869 + (1.058 x 0.86)] (adjx)

× [0.8524 + (0.3208 x 0.35)] (NOxadj)

× [1.2835 – (0.0567 x 2.0)] (Slipadj)

× [0.9636 + (0.0455 x 1.0)] (Sulfuradj)

× [15.16 – (0.03937 x 650) + (0.0000274 x 6502)] (Temperatureadj)

= 2,400 ft3

The catalyst and SCR cross-sectional areas using Equations 2.28 and 2.29 are:

2

265

min

60 16

255,000 ft

s

s

ft

acfm Acatalyst 



























2 2

A 1.15 265 ft 305 ft SCR  
Chapter 2 – Selective Catalytic Reduction

2-90

The length and width of the reactor using Equation 2.30 is:

l w 305 17.5 ft 1 2  

The first estimate of the number of catalyst layers using Equation 2.31 is:

3.0

3.1 265

2,400





nlayer 

Rounding this value gives, nlayer = 3.

Checking the actual catalyst height using Equation 2.32:

1 4.0

3 265

2,400

2

3

 





ft

ft hlayer

This value is within the design height limits of 2.5–5 feet.

The total number of catalyst layers is determined by Equation 2.33 with one empty 

catalyst layer:

ntotal = 3 + 1 = 4

The SCR height, excluding the outlet duct and hoppers using Equation 2.34 is:

hSCR = 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 NO2, 46.01g/mole. For an SRF of 1.05, the reagent mass flow rate is given by 

Equation 2.35:

hr

lb

mole

lb

mole

lb

hr

MMBtu

MMBtu

lb

m

N Ox

reagent 64

46.01

0.35 550 0.857 1.05 17.03



   

 

The mass flow rate of 29% aqueous ammonia solution is given by Equation 2.36:

hr

lb hr

lb

msol 221

0.29

64

  

The solution volume flow rate can then be calculated from Equation 2.37 where sol is the

density of the 29% aqueous ammonia solution, 56.0 lb/ft3at 60oF, and the conversion factor is 

7.481 gal/ft3:
Chapter 2 – Selective Catalytic Reduction

2-91 gph

ft

lb

ft

gal

hr

lb

qsol 30

56.0

221 7.481

3

3







The 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 requirement is given by Equation 2.38:

  gal

day

hr Vol gph days tank 9,924 24 30 14 















  

The onsite storage requirement for ammonia solution is approximately 10,000 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 <500 ft above sea level (14.7 

psia):

1.0

14.7

14.7

 

psia

psia ELEVF

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 2.47. The SCRcost, RPC, APHC and BPC must 

be calculated individually using equations 2.48, 2.49, 2.50 and 2.51, respectively. These 

calculations are shown below.

1.00 1 1 $10,926,654 550 0.1

80

85.7% 270,000

0.2 0.9 2

   











   











 

hr

MMBtu SCRCost 490,000 0.35 550 0.857 1 $1,756,169

0.2 5

  



























  

hr

MMBtu

MMBtu

lb RPC 69,000 0.1 550 / 1 0 1 $0

0.7 8 APHC    MMBtu hr    
Chapter 2 – Selective Catalytic Reduction

2-92 460,000 0.1 550 / 1 1 1 $2,475,762 0.4 2 BPC    MBtu hr    

TCI 1.3$10,926,654$1,756,169$0$2,475,762  $19,706,159

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 82%, resulting in a total capacity factor of 82%. As discussed in section 2.6.1, 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 2.57:

Annual MaintenanceCost  0.005$19,706,159  $98,531/ yr

An estimate for power consumption is given by Equation 2.61:

k W

MW

k W

hr

MMBtu

MMBtu

MWh P (1 1) 308

100

0.1 550 1,000 0.56 0.4 3   

















































Electricity cost can then be estimated from Equation 2.62:

yr kWh yr

hr AnnualElec tricity Cost k W $0.067 $148,430 0.82 8,760

 308    

Reagent cost is estimated using Equation 2.58:

yr lb yr

hr

hr

lb Annual Reagent Cost $ $218,370

 64 8,760 0.820.475 

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 

methodology 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, Rlayer = 3, and the total 

catalyst replacement cost is given by:

$125,231

3

$160 2,400 3

3







ft

ft

Catalyst Replacement Cost
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 hyear is 8,760 hr/yr, and Y is given by:

2.7 3

8,760

24,000

  

hours

hours Y

And the future worth factor for the catalyst replacement is given by:

 

0.3111

1 0.07 1

1

0.07 3

 

FWF  

The annual catalyst replacement cost can then be estimated using Equation 2.64:

 

yr

Annual Catalyst placement Cost $38,953 Re  0.3111 $125,231 

Alternatively, using catalyst replacement cost methodology 2, the annual catalyst 

replacement cost can be estimated using Equation 2.67:

m yr

ft

hr ft

MMBtu

placement Cost

Annual Catalyst $160 35.3 $127,074

80

85.7

) (0.4) (1)

550 0.1

Re 3

3

3

0.7 1

2.9

































 











     















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:

yr yr yr yr yr

Total Direct AnnualCost $98,531 $148,430 $218,370 $38,953 $504,283    

Alternatively, if using catalyst replacement cost methodology 2, the total direct annual 

cost is:

yr yr yr yr yr

Total Direct AnnualCost $98,531 $148,430 $218,370 $127,074 $592,405    

As discussed in section 2.4.2, property taxes and overhead are both assumed to be zero, 

and insurance costs are assumed to 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 2.69 as:

  yr

day hour

hours

year

days AdministrativeCharges 0.4 $98,531 $3,580 /333 4 $60 0.03 

















  















   

The capital recovery factor, CRF, is defined by Equation 2.71 as:
Chapter 2 – Selective Catalytic Reduction

2-94  

 

0.0806

1 0.07 1

0.07 1 0.07

3 0

3 0



 



CRF 

and the capital recovery is calculated from Equation 2.70:

yr yr

Capital Recovery $19,706,159 $1,588,049 0.0806 

The total indirect annual costs (IDAC) are calculated in Equation 2.68:

yr yr yr

IDAC $3,580 $1,588,049 $1,591,629  

The total annual cost is the sum of the direct annual and indirect annual costs given by 

Equation 2.72. If using catalyst replacement cost methodology 1, the total annual cost is:

yr yr yr

Total AnnualCost $504,283 $1,591,629 $2,095,912  

Alternatively, if using catalyst replacement cost methodology 2, the total annual cost is:

yr yr yr

Total AnnualCost $592,405 $1,591,629 $2,184,034  

The annual cost in terms of NOx removed can be calculated using the total annual cost 

and the tons of NOx removed annually. The annual reduction in NOx emissions is given by 

Equation 2.11:

tons yr

ton

lb

yr

hr

hr

MMBtu

MMBtu hr

lb

NO Removed yr x 593 /

2,000

0.857 550 0.82 8,760

/

0.35

/ 

   



and the cost effectiveness is estimated using Equation 2.73. If using catalyst replacement cost 

methodology 1, the cost effectiveness is:

tons ton

yr Cost of NO Removal x

$3,534

593

$2,095,912

 

Alternatively, if using catalyst replacement cost methodology 2, the cost effectiveness is:
Chapter 2 – Selective Catalytic Reduction

2-95

tons ton

yr Cost of NO Removal x

$3,683

593

$2,184,034

 
Chapter 2 – Selective Catalytic Reduction

2-96

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