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OAR Box 1870 

Prepped by Ollie Stewart 

Document Number: 

26) ll-A-30 

Docket Number: 

A-90-19 

Printed 11/28/2008 7:05:04 AM Sheet 26 of 67 
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A-90-19 

ll-A-30 

JLE*AIRDOCKET 

OAQPS CONTROL COST MANUAL 

Fourth Edition 

EPA 450/3-90-006 

January 1990 

Sjg 

United States Environmental Protection Agency 

Office of Air Quality Planning and Standards 

Research Triangle Park, North Carolina 27711 

"1 Air 
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OAQPS CONTROL COST MANUAL 

Fourth Edition 

EPA 450/3-90-006 

January 1990 

2£ 

United States Enrironmental Protection Agency 

Office of Air Quality Planning and Standards 

Research Triangle Park, North Carolina 27711 
P.4 

This fourth edition of the OA QPS Control Cost Manual was prepared 

by the Emissions Standards Division of the Office of Air Quality Planning 

and Standards, U.S. Environmental Protection Agency, Research Triangle 

Park, NC 27711. Mention of trade names or commercial products is not 

intended to constitute endorsement or recommendation for use. Copies of 

this report are available through the Library Services Office (MD-35), U.S. 

Environmental Protection Agency, Research Triangle Park NC 27711, or 

from the National Technical Information Service, 5285 Port Royal Road, 

Springfield VA 22161. 

Questions and comments should be addressed to the principal author, 

William M. Vatavuk, OAQPS, phone 919-541-5309 (FTS 629-5309). The 

technical editor for the fourth edition is Ginny Moyer, OAQPS. 
P.5 

Chapter 3 

THERMAL and CATALYTIC 

INCINERATORS 

Donald R. van der Vaart 

James J. Spivey 

Research Triangle Institute 

Research Triangle Park, NC 27709 

William M. Vatavuk 

Al Wehe 

Standards Development Branch, OAQPS 

U.S. Environmental Protection Agency 

Research Triangle Park, NC 27711 

November 1989 

Contents 

3.1 Introduction . 3-3 

3.2 Process Description 3-4 

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3.2.1 Thermal Incinerators 3-7 

3.2.1.1 Direct Flame Incinerators 3-9 

3.2.1.2 Recuperative Incinerators 3-10 

3.2.1.3 Regenerative Incinerators 3-11 

3.2.2 Catalytic Incinerators 3-13 

3.2.2.1 Fixed-Bed Catalytic Incinerators 3-15 

3.2.2.2 Fluid-Bed Catalytic Incinerators 3-16 

3.2.3 Other Considerations—(Packaged versus Field-Erected 

Units, Auxiliary Equipment) 3-17 

3.2.3.1 Packaged vs. Field-Erected Units 3-17 

3.2.3.2 Add Gas Scrubbers 3-17 

3.2.3.3 Heat Exchangers (Preheaters and other Waste 

Energy Recovery Units) '3-18 

3.2.3.4 Other Auxiliary Equipment 3-18 

3.2.4 Technology Comparison 3-19 

3.3 General Treatment of Material and Energy Balances 3-21 

3.4 Design Procedures 3-22 

3.4.1 Steps Common to Thermal and Catalytic Units . . ". 3-23 

3.4.2 Steps Specific to Thermal Units 3-28 

3.4.3 Steps Specific to Catalytic Units 3-36 

3.5 Cost Analysis 3-42 

3.5.1 Estimating Total Capital Investment 3-42 

3.5.1.1 Equipment Costs, EC 3-43 

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3.5.1.2 Installation Costs 3-51 

3.5.2 Estimating Total Annual Cost 3-51 

3.5.2.1 Direct Annual Costs 3-51 

3.5.2.2 Indirect Annual Costs 3-57 

3.5.3 Cost Comparison for Example Case 3-58 

3.6 Acknowledgements 3-58 

Appendix 3A - Properties of Selected Compounds . 3-60 

References . 3-64 

3.1 Introduction 

Incineration, like carbon adsorption, is one of the best known methods of 

industrial gas waste disposal. Unlike carbon adsorption, however, incineration is an ultimate disposal method in that the objectionable combustible 

compounds in the waste gas are converted rather than collected. On the 

other hand, carbon adsorption allows recovery of organic compounds which 

may have more value as chemicals than just their heating value. A major 

advantage of incineration is that virtually any gaseous organic stream can 

be incinerated safely and cleanly, provided proper engineering design is 

used. 

The particular application of both thermal and catalytic incineration 

to gaseous waste streams containing volatile organic compounds (VOC) is 

discussed here. The U.S. Environmental Protection Agency defines any 

organic compound to be a VOC unless it is specifically determined not to 

be a VOC. Indeed, a number of organics {e.g., methane) are specified as not 

being VOCs. Although both VOC and non-VOC organic compounds are 

combustible and are therefore important in the design of the incinerator, 

this distinction is important since it is the control of VOCs that is regulated. 

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3.2 Process Description 

Seldom is the waste stream to be combusted a single organic compound. 

Rather, it is common to have a complex mixture of organic compounds. 

This mixture is typically analyzed for carbon, hydrogen, oxygen, and other 

elements; and an empirical formula is developed which represents the mixture. Combustion of such a mixture of organic compounds containing carbon, hydrogen, and oxygen is described by the overall exothermic reaction: 

C,Htf0, + .•K *C03 + |H, 0 (3.1) 

The complete combustion products COj and H50 are relatively innocuous, making indneration an attractive waste disposal method. When chlorinated or sulfur-containing compounds are present in the mixture, the 

products of complete combustion indude the add components HCI or S03 , 

respectively, in addition to HjO and CO3. In general, these streams would 

require removal of the add components by a scrubber unit, which could 

greatly affect the cost of the indneration system. (The sizing and costing of these scrubbers is covered in the "Wet Scrubbers" chapter of this 

Manual) 

The heart of an indnerator system is a combustion chamber in which 

the VOC-containing waste stream is burned. Since the inlet waste gas 

stream temperature is generally much lower than that required for combustion, energy must be supplied to the indnerator to raise the waste gas 

temperature. Seldom, however, is the energy released by the combustion of 

the total organics (VOCs and others) in the waste gas stream suffident to 

raise its own temperature to the desired levels, so that auxiliary fuel (e.g., 

natural gas) must be added. 

The combustion of the waste gases may be accomplished in a thermal 

indnerator or in a catalytic incinerator. In the catalytic incinerator a catalyst is used to increase the rate of the combustion reaction, allowing the 

combustion to occur at lower temperatures. Because the catalytic process 

operates at a lower temperature than the thermal process, less auxiliary 

fuel may be required in the catalytic process to preheat the waste gas. 

Auxiliary fuel requirements may also be decreased, and energy effidency 

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improved, by providing heat exchange between selected inlet streams and 

the effluent stream. The effluent stream containing the products of combustion, along with any inerts that may have been present in or added to the 

inlet streams, can be used to preheat the incoming waste stream, auxiliary 

air, or both via a "primary", or recuperative, heat exchanger. It is useful 

to define the fractional energy recovery by the preheater, or primary, heat 

exchanger as follows: 

_ Energy actually recovered from flue gas 

EncrSy =T7—- ~ l iT^75 u~ 3 '2 

n Maximum energy recoverable if flue gas approaches 

lowest temperature available to heat exchanger 

The energy actually recovered, the numerator of Equation 3.2, is the increase in sensible heat of the gas, i.e., waste gas or waste gas plus dilution 

air, being heated. The maximum energy recoverable would be the decrease 

in sensible heat of the flue gas, if it were cooled to the temperature of 

the incoming waste gas. While this maximum energy recovery would be 

attained only with a very large heat exchanger, the concept of fractional 

energy recovery is useful in expressing the extent of the improvement in 

energy effidency using a "primary" heat exchanger. 

Energy effidency can be further improved by pladng another ("secondary" ) exchanger downstream of the primary exchanger to recover additional energy from the effluent stream [e.g., to generate low pressure process 

steam or hot water). However, secondary energy recovery is generally not 

used, unless there is a spedfic on-site use for the steam or hot water. 

The majority of industrial gases that contain VOCs are dilute mixtures 

of combustible gases in air. In some applications, such as air oxidation 

processes, the waste gas stream is very defident in oxygen. Depending on 

the oxygen content of the waste stream, auxiliary air may be required to 

combust the total organic content of the waste gas as well as any auxiliary 

fuel that has been used. 

The concentration of combustible gas in the waste gas stream plays 

an integral role in the design and operation of an indnerator. From a 

cost standpoint, the amount of air in excess of the stoichiometric amounts 

should be minimized. For safety reasons, however, any mixture within the 

flammability limits, on either the fuel-rich or fuel-lean side of the stoichiometric mixture, presents an unacceptable fire hazard as a feed stream to 

the indnerator. The lower, or fuel-lean, explosive limit (LEL) of a given 

3-5 
organic compound defines the minimum concentration of that compound 

in air that can produce more energy than is needed to raise its own temperature to the ignition point (i.e., ignite). Similarly, the upper, or fuel-rich, 

explosive limit (UEL) represents the highest concentration of the organic 

in air that is ignitable. In the latter case, air is limiting the reaction. Both 

the LEL and the UEL are measured at ambient conditions. Empirically, it 

has been found that mixtures of hydrocarbons in air at thdr LEL have a 

heating value of approximately 50 Btu/scf. 

Since the majority of industrial waste gases that contain VOCs are dilute mixtures of combustible gases in air, thdr heating value is low and 

thdr oxygen content exceeds that required to combust both the waste organics (VOCs and others) and the auxiliary fuel. If a waste gas above 50 

percent LEL (about 25 Btu/scf) is encountered, it must be diluted to satisfy fire insurance regulations. Generally, the streams are brought to bdow 

25 percent LEL, although concentrations from 25 percent to 50 percent are 

permitted provided the waste stream is continuously monitored by LEL 

monitors. Because air is the usual diluent gas, care must be taken with 

preheating the diluted stream so that it remains bdow about 1200°F. (See 

discussion below on preheating.) A table showing LEL, UEL, and heats of 

combustion for selected organic compounds is given in Appendix 3A. 

The goal of any indneration system is to control the amount of VOCs 

rdeased to the environment. Performance of a control device such as an 

indnerator can be described by a control effidency, defined according to 

the following equation: 

* j. ,«.. . flnlet mass rate VOC - Outlet mass rate VOC1 ,__ ._ .. 

Control EfF.,% = =-r- 7777= x 100 (3.3) 

I Inlet mass rate VOC J 

It is important to note, however, that incomplete combustion of the inlet 

VOCs could result in the formation of other VOCs not originally present. 

For example, the incomplete oxidation of dichloroethane can yield vinyl 

chloride. Both of these compounds are VOCs. The definition given in 

Equation 3.3 would still be meaningful, however, as long as the newly 

formed VOC {e.g., vinyl chloride) is detected. This situation necessitates 

the complete chemical analysis of the inlet and outlet gas streams to confirm 

compliance with State and Federal regulations. 

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Emission Source 

Dilution Air 

Aux Fuel Qaf 

AirrAir Qa 

Combustion 

Chamber 

«». Taste Ges 

Prebeater 

«v. 

Secondary 

Energy Recovery Stack 

Figure 3.1: Thermal Indnerator - General Case 

Performance of an indnerator can also be measured solely by the outlet 

VOC concentration, usually in ppmv. 

There are a number of different indnerator designs. These designs can 

be broadly dassified as thermal systems and catalytic systems. Thermal 

systems may be direct flame indnerators with no energy recovery, flame 

inrinerators with a recuperative heat exchanger, or regenerative systems 

which operate in a cydic mode to achieve high energy recovery. Catalytic 

systems include fixed-bed (packed-bed or monolith) systems and fluid-bed 

systems, both of which provide for energy recovery. The following sections 

discuss design aspects of these systems. 

3.2.1 Thermal Incinerators 

The heart of the thermal incinerator is a nozzle-stabilized flame maintained 

by a combination of auxiliary fuel, waste gas compounds and supplemental air added when necessary (see Figure 3.1). Upon passing through the 

flame, the waste gas is heated from its inlet temperature {e.g., 100°F) to its 

ignition temperature. The ignition temperature varies for different compounds and is usually determined empirically. It is the temperature at 

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which the combustion reaction rate (and consequently the energy production rate) exceeds the rate of heat losses, thereby raising the temperature 

of the gases to some higher value. Thus, any organic/air mixture will ignite 

if its temperature is raised to a suffidently high levd. 

The organic-containing mixture ignites at some temperature between 

the preheat temperature and the reaction temperature. That is, ignition, 

as defined in this section, occurs at some point during the heating of a 

waste stream as it passes through the nozzle-stabilized flame regardless of 

its concentration. The mixture continues to react as it flows through the 

combustion chamber. 

The required levd of VOC control of the waste gas that must be achieved 

within the time that it spends in the thermal combustion chamber dictates 

the reactor temperature. The shorter the residence time, the higher the 

reactor temperature must be. The nominal residence time of the reacting 

waste gas in the combustion chamber is defined as the combustion chamber 

volume divided by the volumetric flow rate of the gas. Most thermal units 

are designed to provide no more than 1 second of residence time to the 

waste gas with typical temperatures of 1,200 to 2,000°F. Once the unit 

is designed and built, the residence time is not easily changed, so that 

the required reaction temperature becomes a function of the particular 

gaseous spedes and the desired level of control. Table 3.1 illustrates the 

variability in (theoretical) reactor temperatures that is required to destroy 

99.99 percent of the inlet mass of various noxious compounds with excess 

air for a 1-second reactor residence time [1]. 

These temperatures cannot be calculated a priori, although indnerator vendors can provide guidelines based on thdr extensive experience. In 

practice, most streams are mixtures of compounds, thereby further complicating the prediction of this temperature. Other studies [2,3,4], which are 

based on actual field test data, show that commerdal indnerators should 

generally be run at 1600°F with a nominal residence time of 0.75 seconds 

to ensure 98% destruction of non-halogenated organics. In some States 

the reactor temperature and residence time of the unit are specified rather 

than attempting to measure actual levels of VOC control. The selected 

temperature must be maintained for the full, selected residence time for 

combustion to be complete. 

These three studies also condude that mixing is a critical factor in determining the destruction effidency. Even though it cannot be measured, 

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P.13 

Table 3.1: Theoretical Reactor Temperatures Required for 99.99 Percent 

Destruction by Thermal Indneration for a 1-Second Residence Time" 

Compound 

acrylonitrile 

allyl chloride 

benzene 

chlorobenzene 

1,2-dichloroethane 

methyl chloride 

toluene 

vinyl chloride 

Temperature, °F 

1,344 

1,276 

1,350 

1,407 

1,368 

1,596 

1,341 

1,369 

'Reference [1] 

mixing is a factor of equal or even greater importance than other parameters, such as temperature. The most feasible and effident way to improve 

the mixing in an indnerator is to adjust it after start-up. The 98% control 

levd discussed in the previous paragraph presumes such an adjustment. 

Ultimatdy, once the unit is built, it is the responsibility of the user to 

operate and maintain the indnerator to insure compliance with applicable 

regulations. 

3.2.1.1 Direct Flame Incinerators 

Many configurations of thermal indnerators exist with the same goal—to 

raise the VOC-containing stream to the desired reaction temperature and 

hold it there for the given reaction time to achieve the required destruction effidency. The simplest example of such a system is the direct flame 

indnerator. With reference to Figure 3.1, the direct flame indnerator is 

comprised only of the combustion chamber. The waste gas preheater and 

the secondary energy recovery heat exchanger are energy recovery devices 

and are not included as part of the direct flame indnerator. 

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P.14 

3.2.1.2 Re operative Incinerators 

Recuperative indnerators have improved energy effidency as a result of 

pladng heat exchangers in the hot outlet gas streams. With reference 

to Figure 3.1, the recuperative indnerator is comprised of the combustion 

chamber, the waste gas preheater, and, if appropriate, the secondary energy 

recovery heat exchanger. 

Primary Energy Recovery (Preheating Inlet Streams) Considerable fuel savings can be realized by using the exit (product) gas to preheat 

the incoming feed stream, combustion air, or both via a heat exchanger, as 

shown in Figure 3.1 in the so-called "recuperative" indnerator. These heat 

exchangers can recover up to 70% of the energy (enthalpy) in the product 

gas. 

The two types of heat exchangers most commonly used are plate-to-plate 

and shell-and-tube. Plate-to-plate exchangers offer high effidency energy 

recovery at lower cost than shell-and-tube designs. Also, because of their 

modular configuration, plate-to-plate units can be built to achieve a variety 

of effidendes. But when gas temperatures exceed 1000°F, shell-and-tube 

exchangers usually have lower purchase costs than plate-to-plate designs. 

Moreover, shell-and-tube exchangers offer better long-term structural reliability than plate-to-plate units.[5] In any case, because most indnerators 

installed are packaged units, the design (and cost) of the recuperative heat 

exchangers have already been incorporated. 

Most heat exchangers are not designed to withstand high temperatures, 

so that most of the energy needed to reach ignition is supplied by the 

combustion of fud in the combustion chamber and only moderate preheat 

temperatures are sought in practice (<1200°F). 

Secondary Energy Recovery (Additional Waste Energy Recovery) It should be noted, however, that at least some of the energy added 

by auxiliary fuel in the traditional thermal units (but not recovered in preheating the feed stream) can still be recovered. Additional heat exchangers 

can be added to provide process heat in the form of low pressure steam or 

hot water for on-site application. Obviously, an in-plant use for such low 

levd energy is needed to realize these savings. 

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P.15 

The need for this higher level of energy recovery will be dependent 

upon the plant site. The additional heat exchanger is often provided by 

the indneration unit vendor. The cost of this additional heat exchanger 

may be estimated via standard heat exchanger correlations and should be 

added to the costs estimated using the cost correlations in this chapter. 

3.2.1.3 Regenerative Incinerators 

A distinction in thermal incinerators can now be made on the basis of this 

limitation in the preheat temperature. The traditional approach to energy 

recovery in thermal units (shown schematically in Figure 3.1) still requires a 

significant amount of auxiliary fud to be burned in the combustion chamber 

when the waste gas heating values are too low to sustain the desired reaction 

temperature at the moderate preheat temperature employed. Additional 

savings can, under these conditions, be realized in units with more complete 

transfer of exit stream energy. This is the concept behind the so-called 

excess-enthalpy or regenerable burner systems. These systems use direct 

contact heat exchangers constructed of a ceramic material that can tolerate 

the high temperatures needed to achieve ignition of the waste stream. 

The operation of the regenerative system is illustrated in Figure 3.2. 

The inlet gas first passes through a hot ceramic bed thereby heating the 

stream (and cooling the bed) to its ignition temperature. If the desired 

temperature is not attainable, a small amount of auxiliary fud is added in 

the combustion chamber. The hot gases then react (releasing energy) in 

the combustion chamber and while passing through another ceramic bed, 

thereby heating it to the combustion chamber outlet temperature. The 

process flows are then switched, now feeding the inlet stream to the hot 

bed. This cyclic process affords very high energy recovery (up to 95%). 

The higher capital costs assodated with these high-performance heat exchangers and combustion chambers may be offset by the increased auxiliary 

fuel savings to make such a system economical. The costs of these regenerative units will be given separatdy in the cost corrdations presented in 

Section 3.5. Regenerative indnerators are not packaged units but are fielderected only. Accordingly, the costs given in Section 3.5 for regenerative 

units are for field-erected units. 

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AaxAlr Mode A 

SUok 

Anx fual 

UodeB 

Aaz /Ml 

Figure 3.2: Regenerable-type Thermal Indnerator 

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Aux Fuel Q»f 

Aux Air Qa 

Prebeat 

Chamber 

« - t 

Catalyst 

Chamber 

\ I Waste Gas 

Prebeater 

*v. 

Emission Source 

Dilution Air 

9f. Secondary 

Energy 

Recovery 

L

-

Stack 

Figure 3.3: Catalytic Indnerator 

3.2. 2 Catalyti c Incinerator s 

Catalytic indnerators employ a bed of active material (catalyst) that facilitates the overall combustion reaction given in Equation 3.1. The catalyst 

has the effect of increasing the reaction rate, enabling conversion at lower 

reaction temperatures than in thermal indnerator units. Nevertheless, the 

waste stream must be preheated to a temperature sufficiently high (usually 

from 300 to 900°F) to initiate the oxidation reactions. The waste stream is 

preheated either directly in a preheater combustion chamber or indirectly 

by heat exchange with the indnerator's effluent or other process heat or 

both (Figure 3.3). The preheated gas stream is then passed over the catalyst bed. The chemical reaction (combustion) between the oxygen in the 

gas stream and the gaseous pollutants takes place at the catalyst surface. 

Catalytic indneration can, in prindple, be used to destroy essentially any 

oxidizable compound in an air stream. However, there are practical limits 

to the types of compounds that can be oxidized, due to the poisoning effect. 

some spedes have on the catalyst. These limits are described below. In 

addition, most configurations require a low heating value of the inlet gas 

and a particulate content which is less than some small value. 

Until recently, the use of catalytic oxidation for control of gaseous pels'13 
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lutants has generally been restricted to organic compounds containing only 

carbon, hydrogen and oxygen. Gases containing compounds with chlorine, 

sulfur, and other atoms that may deactivate the supported noble metal catalysts often used for VOC control were not suitably controlled by catalytic 

oxidation systems. Catalysts now exist, however, that are tolerant of such 

compounds. Most of these catalysts are single or mixed metal oxides, often 

supported by a mechanically strong carrier such as 7-alumina. Perhaps 

most of the development of poison-tolerant catalysts has focused on the 

oxidation of chlorine-containing VOCs. These compounds are widdy used 

as solvents and degreasers and are often the subject of concern in VOC control. Catalysts such as chromia/alumina [6,7], cobalt oxide [8], and copper 

oxide/manganese oxide [9] have been used for oxidation of gases containing 

chlorinated compounds. Platinum-based catalysts are active for oxidation 

of sulfur-containing VOCs, although they are rapidly deactivated by the 

presence of chlorine. Compounds containing atoms such as lead, arsenic,. 

and phosphorous should, in general, be considered poisons for most oxidation catalysts. Neverthdess, thdr concentration may be suffidently low so 

that the rate of deactivation and therefore, the catalyst replacement costs, 

could be low enough to consider catalytic oxidation. 

As was the case for thermal units, it is impossible to predict a priori 

the temperature and residence time (i.e., inverse space vdodty) needed to 

obtain a given levd of conversion of a VOC mixture in a catalytic oxidation 

system. For example, Table 3.2 from Pope et al [8] shows the temperature 

needed for 80% conversion of a number of VOCs over two oxidation catalysts in a specific reactor design. This table shows that the temperature 

required for this levd of conversion of different VOCs on a given catalyst 

and of the same VOC on different catalysts can vary significantly. 

Particulate matter, induding dissolved minerals in aerosols, can rapidly 

blind the pores of catalysts and deactivate them over time. Because essentially all the active surface of the catalyst is contained in rdativdy small 

pores, the particulate matter need not be large to blind the catalyst. No 

general guidelines exist as to particulate concentration and particulate size 

that can be tolerated by catalysts because the pore size and volume of 

catalysts vary greatly. 

The volumetric gas flow rate and the concentration of combustibles in 

the gas flowing to the catalytic indnerator should be constant for optimal 

operation. Large fluctuations in the flow rate will cause the conversion 

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P.19 

Table 3.2: Catalyst Temperatures Required for Oxidizing 80% of Inlet 

VOC to COj, °F for Two Catalysts 

Compound Temperature, °F 

C03O4 Pt - Honeycomb 

acrolein 382 294 

n-butanol 

n-propylamine 

toluene 

n-butyric add 

1,1,1-trichloroethane 

dimethyl sulfide 

413 

460 

476 

517 

661 

— 

440 

489 

373 

451 

>661 

512 

of the VOCs to fluctuate also. Changes in the concentration or type of 

organics in the gas stream can also affect the overall conversion of the 

VOC contaminants. These changes in flow rate, organics concentration, 

and chemical composition are generally the result of upsets in the manufacturing process generating the waste gas stream. It may be uneconomical 

to change the process for tho sake of making the operation of the catalytic 

incinerator feasible. In such cases, thermal indnerators (discussed earlier in 

this chapter) or carbon adsorption (discussed in Chapter 4 of this Manual) 

should be evaluated as alternative control technologies. 

The method of contacting the VOC-containing stream with the catalyst 

serves to distinguish catalytic indneration systems. Both fixed-bed and 

fluid-bed systems are used. 

3.2.2.1 Fixed-Bed Catalytic Incinerators 

Fixed-bed catalytic indnerators may use a monolith catalyst or a packedbed catalyst. Each of these is discussed bdow. 

Monolit h Catalyst Incinerators The most widespread method of contacting the VOC-containing stream with the catalyst is the catalyst mono3-15 
P.20 

lith. In this scheme the catalyst is a porous solid block containing parallel, non-intersecting channels aligned in the direction of the gas flow. 

Monoliths offer the advantages of minimal attrition due to thermal expansion/contraction during startup/shutdown and low overall pressure drop. 

Packed-Bed Catalytic Incinerators A second contacting scheme is a 

simple packed-bed in which catalyst partides are supported dther in a tube 

or in shallow trays through which the gases pass. The first scheme is not 

in widespread use due to its inherently high pressure drop, compared to a 

monolith, and the breaking of catalyst partides due to thermal expansion 

when the confined catalyst bed is heated/cooled during startup/shutdown. 

However, the tray type arrangement, where the catalyst is pelletized is used 

by several industries {e.g., heat-set web-offset printing). Pelletized catalyst 

is advantageous where large amounts of such contaminants as phosphorous 

or silicon compounds are present.[10] 

3.2.2.2 Fluid-Bed Catalytic Incinerators 

A third contacting pattern between the gas and catalyst is a fluid-bed. 

Fluid-beds have the advantage of very high mass transfer rates, although 

the overall pressure drop is somewhat higher than for a monolith. An additional advantage of fluid-beds is a high bed-side heat transfer as compared 

to a normal gas heat transfer coeffident. This higher heat transfer rate to 

heat transfer tubes immersed in the bed allows higher heat release rates 

per unit volume of gas processed and therefore may allow waste gases with 

higher heating values to be processed without exceeding maximum-permissible temperatures in the catalyst bed. In these reactors the gas phase 

temperature rise from gas inlet to gas outlet is low, depending on-the extent of heat transfer through imbedded heat transfer surfaces. The catalyst 

temperatures depend on the rate of reaction occurring at the catalyst surface and the rate of heat exchange between the catalyst and imbedded heat 

transfer surfaces. 

As a general rule, fluid-bed systems are more tolerant of particulates 

in the gas stream than dther fixed-bed or monolithic catalysts. This is 

due to the constant abrasion of the fluidized catalyst pellets, which hdps 

remove these particulates from the exterior of the catalysts in a continuous 

manner. 

3-16 
P.21 

A disadvantage of a fluid-bed is the gradual loss of catalyst by attrition. Attrition-resistant catalysts have been developed to overcome this 

disadvantage.[ll] 

3.2.3 Other Considerations—(Packaged versus FieldErected Units, Auxiliary Equipment) 

3.2.3.1 Packaged vs. Field-Erected Units 

With the exception of regenerative indnerators, the equipment cost correlations included in this chapter are for packaged units only. They are 

not valid for fidd-erected units. For regenerative indnerators, the correlations are valid for fidd-erected units only. Packaged units are units that 

have been shop fabricated and contain all dements necessary for operation, 

except for connection to fadlities at the site, e.g., utilities. The elements 

indude the combustion chamber, preheater, instrumentation, fan, and the 

necessary structural sted, piping, and dectrical equipment. This equipment is assembled and mounted on a "skid" to facilitate installation on a 

foundation at the plant site. Tie-in to the local emission source is not part 

of the packaged unit. Units are usually sized to handle flow rates of <20,000 

scfm, but can be built to accommodate flow rates up to 50,000 scfm. The 

cost corrdations in this chapter are valid to 50,000 scfm for packaged units, 

except for fluid-bed units which are valid to 25,000 scfm. 

Conversdy, field-erected units may be built to any desired size. The 

combustion chamber, preheater, and other equipment items are designed 

and fabricated individually, and assembled at the site. However, both the 

equipment and installation costs of fidd-erected units are typically higher 

than those for equivalent-sized packaged units because the factors that improve effidency of shop-fabrication, such as uniform working environment, 

availability of tools and equipment, and more effident work scheduling, are 

generally not available in the field. 

3.2.3.2 Acid Gas Scrubbers 

The final outlet stream of any indneration system may contain certain pollutants that must be removed. The combustion of sulfur-containing com3-17 
P.22 

pounds results in S03 , while chlorinated compounds yidd Clj and HCI in 

the product stream. These add gases must be removed from the gas stream 

if they are present at significant concentrations (regulations for limits on 

these gases vary from state to state). This removal can be effected in, for 

instance, a packed-bed vertical scrubber in which the flue gas is contacted 

with a caustic scrubbing liquid. For fluid-bed catalytic reactors, venturi 

scrubbers are often used because they provide for particulate removal as 

well as add gas scrubbing. In most cases adding a scrubber significantly 

increases the cost of the indneration unit, sometimes by a factor of two. 

Costing of scrubbers is discussed in the "Wet Scrubbers" Chapter of this 

Manual. 

If chlorinated VOCs are present in the waste gas, heat exchangers may 

require spedal materials of construction. This added expense is not induded in the costing procedures outlined in this chapter. 

3.2.3.3 Heat Exchangers (Preheaters and other Waste Energy 

Recovery Units) 

For the thermal and catalytic units having some degree of energy recovery, 

the cost of the primary heat exchanger is induded in the cost, and its design is usually done by the indneration unit vendor. The cost corrdations 

presented in this chapter indude units both with and without energy recovery. Secondary energy recovery, if desired, requires an additional heat 

exchanger, which is also often provided by the incineration unit vendor. 

Costing procedures for secondary energy recovery are not induded in this 

chapter. 

3.2.3.4 Other Auxiliary Equipment 

Additional auxiliary equipment such as hoods, ductwork, precoolers, cydones, fans, motors, and stacks are addressed separatdy in other chapters 

of this Manual. 

3-18 
P.23 

3.2.4 Technology Comparison 

Both the thermal and catalytic indneration systems are designed to provide 

VOC control through combustion at a level in compliance with applicable 

state and federal requirements. Given the wide range of options available, 

however, it is obvious that not every indnerator will fulfill these requirements at the same cost. This section presents a first step toward dedding 

how best to deal with VOC emission abatement using indnerators considering some qualitative factors pertinent to the types of indnerators described 

in this chapter. It is the intent of the remainder of Chapter 3 to provide a 

method by which the cost of VOC control for a particular application can 

be calculated. 

A summary of the principal types of indnerators is presented in Table 3.3. From the earlier discussions, the following factors relating to the 

presence of contaminants should be considered by potential users [12]: 

e The fouling of the catalyst in a catalytic system is a possibility. Poisons to the system include heavy metals, phosphorous, sulfur and 

most halogens, although catalysts have been devdoped that are chlorine resistant. 

e The possibility of process upsets that could release any of the above 

poisons or cause fluctuations in the heating value to the indnerator 

would favor a thermal system. 

e Except for the No.2 grade, fud oil should not be considered as auxiliary fud to a catalytic system due to the sulfur and vanadium it may 

contain.(10] 

All of the above factors would serve to increase the operating expense of a 

catalytic unit through replacement costs of the catalyst. An additional factor relates to rdative energy effidency of the various types of indnerators: 

e Thermal units generally require more auxiliary fud than catalytic 

units and operate at temperatures that are roughly 1000° F higher. 

This difference in fud requirement increases as the heating value of 

the waste stream decreases. 

3-19 
P.24 

Table 3.3: Prindpal VOC Indnerator Technologies 

Thermal Systems 

e Direct Flame Indnerator 

e Recuperative Indnerator (Direct Flame with Recuperative Heat Exchanger) 

e Regenerative Indnerator Operating in a Cyclic Mode 

Catalytic Systems 

e Fixed-Bed 

- Monolith 

- Packed-Bed 

• Fluid-Bed 

3-20 
P.25 

In general, a trade-off exists between the higher capital costs of catalytic 

indnerators and the higher operating costs of thermal indnerators. This 

difference will be illustrated by a design example presented in Section 3.4 

which treats both technologies. 

3.3 General Treatment of Material and Energy Balances 

In the sizing and costing of the indnerator and the calculation of the auxiliary fuel requirements, it is necessary to make material and energy balances 

around the entire indnerator unit and around certain parts of the unit, such 

as the combustion chamber or the preheater. This section presents a general 

approach to making these balances. 

These balances are based on the law of conservation of mass and energy. 

They can be stated in general equation form as 

In — Out + Generation — Accumulation (3*4) 

Because the indneration process is a steady-state process, the accumulation 

term is zero and the equation becomes 

In — Out + Generation = 0 

For mcuj balances it is useful to restrict the balances to be made on the 

mass of each atomic spedes so that for mass balances the generation term 

becomes zero. However, because the combustion reaction liberates energy, 

the energy balances around equipment where combustion takes place would 

indude a generation term. Thus, the simplified equations are 

In — Out ss 0 , for steady-state mass balances (3.5) 

In — Out + Generation — 0 , for steady-state energy balances (3.6) 

3-21 
P.26 

For the indneration process the two terms In and Out are generally 

mass terms (for a mass balance) of the form 

PQ, 

where 

p — density (mass per unit volume) 

Q — volumetric flow rate (volume per unit time) 

or sensible heat terms (for an energy balance) of the form, 

PQC,(T-T„,) 

where 

Cp = heat capadty 

T =s temperature. 

The reference temperature, T„/, is often taken to be zero or the temperature of a convenient stream, e.g., the inlet gas stream, in whatever 

units T is in, so the T„/ term may not appear in the equations. When the 

reference temperature is taken as zero, the sensible heat terms become 

PQCJ. 

Energy losses, H&, are also part of the Out term and, for the indnerator 

process, are taken here to be 10% of the total energy tnpui to the indnerator. 

For the indneration process, the generation term for energy balances 

accounts for the energy released through the combustion reactions. This 

term is generally of the form 

PSC-AA,) •* 

where 

(—A&e) = heat of combustion. 

3.4 Design Procedures 

The following procedure is designed to provide parameters for use in devdoping a study cost estimate (accuracy ± 30%). The prindpal parameters 

3-22 
P.27 

of interest are 

e flue gas flow rate, upon which all the equipment cost correlations are 

based. 

e auxiliary fuel requirement, which is important in estimating annual 

operating costs. 

For applications which involve control of waste gas streams that are 

dilute mixtures of VOCs in air (>20% oxygen in the waste gas stream), 

the flue gas flow rate is greater than the inlet waste gas flow rate by the 

amount of auxiliary fuel and the increase in the moles of gas as a result 

of the combustion reaction. Because these two factors usually cause only 

small increases in flow rate, a number of simplifying assumptions can be 

made in the design calculations. For applications where diluent air must be 

used to adjust the combustible concentration in the waste gas to 25% LEL 

and where auxiliary fud and auxiliary combustion air are needed, more 

complete mass and energy balances must be made. 

The design procedure illustrated bdow is for waste gas streams that are 

dilute mixtures of VOCs in air (>20% oxygen in the waste gas stream). In 

this discussion the design procedure will be illustrated by a sample problem 

that will be solved step-by-step. 

3.4.1 Steps Common to Thermal and Catalytic Units 

Step 1 • Establish design specifications The first step in the design 

procedure is to determine the spedfications of the indnerator and the waste 

gas to be processed. The following parameters of the waste gas stream at 

the emission source must be available: 

• Volumetric flow rate, scfm—Standard conditions are normally 77°F 

and 1 atm. pressure. 

• Temperature 

t Oxygen content 

e Chemical composition of the combustibles 

3-23 
P.28 

e Inerts content 

e Heating value—In some cases the heating value may act as a surrogate 

for the chemical composition of the combustibles. This is particularly 

true for dilute mixtures ot combustibles in air. 

e Particulate content—The particulate content is important if catalytic 

indnerators are to be costed. An upstream filter may suffice if particulate content is too high. Fluid-bed catalytic indnerators can tolerate 

higher particulate contents than fixed-bed catalytic indnerators. 

The following parameters must be spedfied for the incinerator: 

e Desired control effidency - This effidency should be based on requirements dictated by relevant state and federal regulations. 

e Combustion chamber outlet temperature - This temperature may 

also be based on requirements of a regulation or on recommendations 

devdoped during regulatory devdopment. 

• Derired percent energy recovery - The desired percent energy recovery should be the result of a process optimization in which costs of 

incinerators with several different levels of energy recovery are estimated and the minimum cost design sdected. The tradeoff is between 

the capital cost of the energy recovery equipment and the operating 

(fud) costs. 

Spedfications for the sample problem are given in Table 3.4. 

Step 2 • Verify that the oxygen content of the waste gas exceeds 

20% There must be suffident oxygen in the waste gas to support the 

combustion of the waste organics (induding VOCs) and the auxiliary fud, 

if auxiliary fud is needed. It may be necessary to add auxiliary air if the 

oxygen content is less than about 20%. This example is based on streams 

that contain >20% oxygen, as shown bdow: 

Air Content, Vol. % - 100.0 - ~ x 100 - ^ ? x 100 (3.7) 

10s 10* 

- 99.8% 

Oxygen Content,% = Air Content x 0.209 (3.8) 

= 20.86% 

3-24 
P.29 

Table 3.4: Spedfication of Sample Problem 

Variable Value 

Preheater Inlet Waste Gas Vol Flow Rate, Q„,, scfm 20,000 

Preheater Inlet Waste Gas Temp., T„<, °F 100 

Composition 

Benzene Content, ppmv 1000 

Methyl Chloride Content, ppmv 1000 

Air Content Balance 

Particulate Content Negligible 

Moisture Content Negligible 

Desired Control Effidency,% 98 

Desired Percent Energy Recovery, HR,% 70 

Step 3 • Calculate the LEL and the Percent of the LEL of the 

gas mixture Note: If the waste stream contains a significant amount of 

inerts in addition to the nitrogen assodated with the oxygen in air, the 

calculation of LEL (and UEL) loses meaning since the LEL (and UEL) is 

measured in mixtures of organic with air only. A complete chemical analysis 

is necessary to complete the design procedure in such a case. 

The example chosen here is typical, in that there is more than one VOC 

component in the gas stream. An approximate method to calculate the 

LEL of a mixture of compounds, LEI*^, is given by Grdecki [13] as 

LEU* IjSt&t'dxLEL^ (3.9) 

where 

LEL; 

n — 

volume fraction of combustible component i 

lower explosive limits of combustible component 

; (ppmv) 

number of combustible components in mixture 

3-25 
P.30 

For the example case, 

J^Xi = (1,000 + 1,000) xl0~a 

t a l 

s= 2,000 xlO - 6 

»• 

From standard references [14] or from Appendix 3A, 

LELB« = 14,000 ppmv for benzene 

LELMC = 82,500 ppmv for methyl chloride 

(3.10) 

LEU 

1,000 

+ 

1,000 1 - 1 

2,000 x 14,000 2,000 x 82,500 

23,938 ppmv 

(3.11) 

% LEL 

total combustible cone, in mixture 

2,000 

23,938 

LEL ^ 

x 100 s= 8.4% 

x 100 (3.12) 

'(3.13) 

The percent LEL of the mixture is therefore 8.4%. Because this is well 

below 25%, no dilution air is needed in this example. If the mixture had 

been above 25% LEL, suffident dilution air would have been needed to bring 

the concentration of the mixture to less than 25% to satisfy fire insurance 

regulations. 

Step 4 • Calculate the volumetric heat of combustion ofthe waste 

gas stream, (-A/O , Btu/scf The energy content of the gas stream, 

expressed in terms of the heat of combustion, is calculated as follows: 

(-AAe.) - B-A/O* ; (3.14) 

3-26 
P.31 

where 

(-A/ic.) — beat of combustion of the waste stream (Btu/scf) 

(-A/ieJ = volumetric heat of combustion of component i at 

25°C(Btu/scf) 

Xj ss volume fraction of component i in the waste ga? 

n = number of combustible components in the waste 

gas 

The heat of combustion that should be used in these calculations is the 

"lower" heat of combustion, i.e., with gaseous water, rather than liquid 

water, as a reaction product since water leaves the indnerator in the vapor 

state. From Appendix 3A or standard references [14,15] with appropriate 

conversion of units, the volumetric heat of combustion at 25°C for the two 

components is calculated to be as follows: 

(-A/teg.) = 3,475 Btu/scf for benzene 

(-AA^ ) = 705 Btu/scf for methyl chloride 

The compositions spedfied earlier as ppmv are converted to volume 

fractions as follows: 

*B« — 1»000 ppmv xlO"6 ss 10"3 for benzene 

*MC - 1,000 ppmv xlO"* = IO"3 for methyl chloride 

Using these values of heat of combustion and composition, the heat of 

combustion of the waste gas stream per standard cubic foot of incoming gas 

is 

(-AA*.) ss (3,475) (l0"3)+(705) (lO-3) (3.15) 

s= 4.18 Btu/scf 

Assuming the waste gas is prindpally air, with a molecular weight of 

28.97 and a corresponding dendty of 0.0739 lb/scf, the heat of combustion 

per pound of incoming waste gas is 

(-A/ 0 = 56.6 Btu/lb 

3-27 
P.32 

The negative heat of combustion, by convention, denotes an exothermic 

reaction. Also by convention, if one refers to heat of reaction rather than 

heat of combustion, then a positive value denotes an exothermic reaction. 

Empirically, it has been found that 50 Btu/scf roughly corresponds to 

the LEL of organic/air mixtures. Insurance codes require a value bdow 25% 

LEL, which corresponds to about 13 Btu/scf. However, if LEL sensors and 

monitors are installed, one can indnerate a waste gas with a combustible 

organic content between 25 and 50% LEL, which corresponds to 13 to 25 

Btu/scf. 

For catalytic applications the heat of combustion must normally be less 

than 10 Btu/scf (for VOCs in air) to avoid excessivdy high temperatures 

in the catalyst bed. This is, of course, only an approximate guideline and 

may vary from system to system. 

After Step 4, determination of (-AA^), the design procedure for thermal and catalytic indnerators is discussed separatdy, beginning with Step 

5 for each type of indnerator. 

3.4.2 Steps Specific to Thermal Units 

Figure 3.1 shows a generic thermal indnerator with the appropriate streams 

labeled. 

Step 5t • Establish the temperature at which the incinerator will 

operate As mentioned in Section 3.2.1, both the reactor temperature and 

residence time of the waste gas in the reactor determine the levd of VOC 

destruction. In general, state and local regulations specify the required levd 

of destruction that the customer must meet. In this example a destruction 

effidency of 98 percent is spedfied. Studies by Mascone [2,3,4] show that 

this destruction effidency can be met in a thermal incinerator operated 

at a temperature, T/{, of 1,600°F and a residence time of 0.75 second. 

(Note: This higher effidency levd is the minimum achievable by any new 

properly designed and operated indnerator. Many indnerators can achieve 

destruction effidendes of 99% or higher.) 

3-28 
P.33 

Step 6t • Calculate the waste gas temperature at the exit of the 

preheater The extent of heat exchange to be carried out in the preheater 

is the result of a technical and economic optimization procedure that is not 

illustrated in this example. As the VOC stream temperature leaving the 

heat exchanger, Tw., increases, the auxiliary fuel requirement decreases, 

but at the expense of a larger heat exchanger. However, there are several 

important Hmits on TWa. First, Tw. must not be dose to the ignition temperature of the organic-containing gas to prevent damaging temperature 

excursions inside the heat exchanger should the gas ignite.. Second, for 

gases containing halogens, sulfur, and phosphorous (or other arid-forming 

atoms), the flue gas temperature after the heat exchanger,.T/#, must not 

drop below the add dew. Both limitations limit the amount of heat exchange and thus the maximum value of T„.. The calculation of the acid 

dew is not simple. It is recommended that vendor guidance be sought to 

ensure that the dew is not reached. Condensation of add gases will result 

in corrosion of many of the metals used in heat exchangers. As an example, 

fud sulfur contents of 1 to 2 percent can give add dew points of about 200 

to 270° F. Increasing the sulfur content to 4 percent can raise the dew to 

about 290° F. Chlorine and phosphorous have a much smaller effect on add 

dew elevation. 

With the following assumptions, one can estimate T„. using equation 

3.2, the definition of fractional energy recovery for a heat exchanger. 

e The fractional energy recovery is spedfied. 

e The amount of auxiliary fuel, Q«/, and auxiliary combustion air, Qa, 

are small relative to the waste'gas, Qw, so that the mass flow rates 

of gases, pwQ* and P/Q/, on both sides of the preheater are approximatdy the same, or 

PwQw * PtQt 

• The heat capadties of the gases on both sides of the preheater are 

approximatdy the same, regardless of composition. This is true for 

waste streams which are dilute mixtures of organics in air, the properties of the streams changing only slightly on combustion. 

e The mean heat capadties above the reference temperature of the gases 

on both sides of the preheater are approximatdy the same regardless 

of temperature. 

3-29 
P.34 

With these assumptions, the equation for fractional energy recovery for a 

heat exchanger becomes 

T - T Fractional Energy Recovery = -£r——- (3.16) 

T/< - Tm 

For this example with a fractional energy recovery of 0.70, an indnerator 

operating temperature, T/<, of 1,600°F, and a waste gas inlet temperature, 

T^, of 100°F, the waste gas temperature at the exit of the preheater becomes 

T«. = 1,150°F 

The temperature of the exhaust gas, T/a, can be determined by an 

energy balance on the preheater, which, with the same assumptions as 

used in deriving Equation 3.16 regarding the mass flow rates and average 

heat capadties of the gases involved, results in the following equation: 

Th - Tf. * T„. - T^ 

i.e., the temperature rise in the waste gas is approximately equal to the 

temperature decrease in the flue gas with which it is exchanged. For this 

example, this results in 

Tu = 550°F 

This value of T/s should be well above the add dew of the flue gas stream. 

It should be remembered that Tw. should be well bdow the ignition temperature of the VOC stream to prevent unwanted temperature excurdons 

in the preheater. This must be verified even if the stream is well below the 

LEL because flammability limits can be expanded by rairing the reactant 

stream temperature. A suffidently high preheat temperature, Tw#, could 

initiate reaction (with heat rdease) in the preheater. This would ordinarily 

be detrimental to the materials of construction in the heat exchanger. The 

one exception is the thermal indnerator of the regenerable type described 

3-30 
P.35 

•in Section 3.2 The 95 percent energy recovery obtainable in regenerable 

systems would result in this example in a Tw# of 1,525°F. The significant 

reaction rate that would occur at this temperature in the ceramic packing 

of the heat exchanger/reactor is by design. 

Step 7t - Calculate the auxiliary fuel requirement, Qa / Auxiliary 

fuel will almost invariably be needed for startup of the unit. However, at 

steady state, if the energy released by combustion of the organics present in 

the waste stream is suffident to maintain the reactor temperature (1,600°F 

in the example), only a small amount of auxiliary fuel (< 5% of the total 

energy input) is needed to stabilize the flame. In most cases, however, more 

fuel than just this stabilizing fuel will be required to maintain the reactor 

temperature. 

With the following assumptions, one can estimate Q„/ using a mass and 

energy balance around the combustion chamber and following the prindples 

discussed in Section 3.3, with reference to Figure 3.1. 

• The reference temperature, T,,/, is taken as the inlet temperature of 

the auxiliary fud, T«/. 

e No auxiliary air, Qa, is required. 

e Energy losses, EL, are assumed to be 10% of the total energy input to 

the indnerator above ambient conditions.[16,17] Thus, if the reference 

temperature is near ambient conditions, 

HL^H.\phqhC^u(T}i-T^) (3.17) 

• The heat capadties of the waste gases entering and leaving the combustion chamber are approximatdy the same, regardless of composition. This is true for waste streams which are dilute mixtures of 

organics in air, the properties of the streams changing only slightly 

on combustion. 

e The mean heat capadties above the reference temperature of the 

waste gases entering and leaving the combustion chamber are approximatdy the same regardless of temperature. Thus the mean heat 

capadty for the waste gas stream entering or leaving the combustion 

chamber should be evaluated at the average of TWa and T/ r For air 

3-31 
P.36 

this assumption introduces an error of, at most, 5% over the temperatures of interest. 

With these assumptions, the mass and energy balance around the combustion chamber reduces to the following equation: 

„ 0 p.;<?..[gp-^(i-ir/, - r.. - o.ir,.,) - (-A/Q I 

'•'«•'= (-So-i.iWJi-JW— "

 (348) 

Input data for this equation are summarized below: 

e The waste stream is essentially air so that 

Pw. = pwi = 0.0739 lb/scf, air at 77"F, 1 atm. 

C^ .r = 0.255 Btu/lb °F, the mean heat capadty of air between 

77°F and 1,375°F (the average temperature of the waste gas entering and 

leaving the combustion chamber) 

• Other input data to Equation 3.18 indude 

Qw. = Qm = 20,000 scfm 

(-A&e.,) - 21,502 Btu/lb, for methane 

7/,,/ ss T9f s= 77* F, assume ambient conditions 

pMf ss 0.0408 lb/ft3, methane at 77°F, 1 atm. 

Tfi ss 1,600°F, from Step 5t 

T*. = 1,150°F, from Step 6t 

(-AAc^) ss 56.6 Btu/lb, from Step 4 

Substituting the above values into Equation 3.18 results in 

Qmf s= 167 scfm 

The values of the parameters in the energy balance are summarized in 

Table 3.5. 

3-32 
P.37 

Table 3.5: Summary of Example Problem Variable Valuation 

T„, == 77°F 

Pj* Qii Qpmji T,> 

Stream Subscript,; lb/scf scfm Btu/lb °F °F 

IN - Sensible Heat 

Auxiliary Air a na* na' na* na* 

Auxiliary Fud af 0.0408 167 " 77 

Waste Gas w9 0.0739 20,000 0.255 1,150 

OUT - Sensible Heat 

Waste Stream f{ 0.0739 20,167 0.255 1,600 

(-AA,), waste gas = 56.6 Bto/lb 

(-AA,). auxiliary fuel = 21,502 Btu/lb 

'Not applicable. 

"Not used because reference temperature is taken equal to auxiliary fuel temperature. 

It is instructive to examine the magnitude of the various terms in the 

energy balance around the combustor for the sample problem. This is done 

in Table 3.6. The energy balance shown does not quite add to zero due 

to round-off error and simplifying assumptions. Table 3.6 shows that the 

largest inlet term is the sennble heat of the incoming waste gas. The heat of 

combustion of the organics contained in the waste gas stream is somewhat 

smaller than that of the auxiliary methane because of the rdativdy small 

amount of organics in the waste gas stream. The largest term in the outlet 

stream is the sensible heat of the outgoing waste stream. The overall energy 

losses are based on an assumption, but are rdativdy small. Because the 

sensible heat contents of the entering and leaving waste stream are so large, 

it is apparent that energy recovery is an important factor in achieving 

energy effidency. In fact, with zero energy recovery in the sample problem, 

the auxiliary fud requirements would be 605 scfm, about four times the 

energy requirements based on 70% energy recovery. 

3-33 
P.38 

Table 3.6: Terms in Energy Balance around Combustor - Example Problem 

IN - OUT + GENERATION = 0 

Stream 

IN - Sensible Heat, pjQjC^Ti - T„f) 

Auxiliary Air 

Waste Gas 

OUT - Sensible Heat, pjQjC^Ti - T„t) 

Waste Stream 

OUT - Losses 

10% of total energy input 

GENERATION -

Heat of Combustion, PjQ^—Ahq) 

Waste Gas 

Auxiliary Fud 

Subscript, j 

a 

fi 

af 

Value, 

Btu/min 

0 

404,403 

578,796 

57,880 

83,655 

146,506 

3-34 
P.39 

Step 8t - Verify that the auxiliary fuel requirement is sufficient to 

stabilize the burner flame Only a small amount of auxiliary fuel (< 5% 

of the total energy input) is needed to stabilize the burner flame. In general, 

more fud than just this stabilizing fud will be required to maintain the 

reactor temperature. It is wise to verify that the auxiliary fuel requirement 

calculated in Step 7t is suffident for stabilization. If it is insuffident, then a 

minimum amount of auxiliary fuel must be used, and the amount of energy 

recovery spedfied earlier must be reduced to avoid exceeding the specified 

temperature at which the indnerator will operate (Step 5t). 

This check is made by calculating 5% of the total energy input to the 

indnerator and comparing it with the auxiliary fuel energy input. The total 

energy input is given as follows: 

Total Energy Input = PuQifi^Tu - T„f) (3.19) 

Auxiliary Fud Energy Input ss p./Q./(-Ane#/) (3.20) 

The auxiliary fud used in the design, Q„/, should be the larger of 5% 

of the total energy input (28,900 Btu/min.) and the auxiliary fuel energy 

input (146,500 Btu/min.). The auxiliary fud used easily meets this criterion. 

Step 9t • Calculate the total volumetric flow rate of gas through 

the incinerator, Q/, The total volumetric flow rate of gas leaving the 

indnerator is referred to as the flue gas flow rate, Qj{1 and is the gas rate on 

which the indnerator sizing and cost correlations are based. The flue gas 

flow rate measured at the standard conditions of 77°F and 1 atmosphere, 

where the increase in volumetric throughput due to an increase in the number of moles of gas as a result of combustion is neglected, is the sum of the 

inlet streams to the indnerator. 

Qh = <?.. + $• + <?./ 

= 20,000 + 0 + 167 

ss 20,167 scfm 

This result conforms with the assumptions stated in Step 6t, i.e., the mass 

3-35 
P.40 

(and volume) flow rates on both sides of the preheater are approximately 

equal. 

3.4.3 Steps Specific to Catalytic Units 

Figure 3.3 shows a generic catalytic indnerator with the appropriate streams 

labeled. The approach used in the calculations on the catalytic indnerator is somewhat different than that used in the thermal indnerator. This 

difference arises because of additional constraints which are placed on the 

catalytic indnerator. These constraints are as follows: 

e The desired catalyst bed outlet temperature is typically 700 to 900 °F. 

The maximum temperature to which the catalyst bed can be exposed 

continuously is limited to about 1,200 "F. Therefore, the combustible 

content of the waste gas is limited, and the amount of heat exchange 

that occurs in the primary heat exchanger may be limited. 

e The inlet temperature to the catalyst bed itself must be above the 

catalytic ignition temperature required to give the derired destruction 

effidency in the indnerator. Therefore, the combustible content of the 

waste gas is further limited to that which, when combusted, will raise 

the temperature in the catalyst bed no more than the AT between 

the required reactor bed inlet temperature, and the dedred reactor 

bed outlet temperature. 

e Auxiliary fud, in combination with the preheat from the primary 

heat exchanger, is used y preheat the waste gas to the reactor inlet temperature. A mi; aum amount of auxiliary fud (< 5% of the 

total energy input) must be used to stabilize the flame in the preheat combustion chamber. This has the effect of farther limiting the 

combustible content of the waste gas stream and the amount of heat 

exchange permisnble in the primary heat exchanger. 

The steps outlined bdow represent one approach to recognizing these 

constraints and incorporating them into the calculation procedures. 

Step 5c • Establish the desired outlet temperature of the catalyst bed, T/{

 The energy rdeased by the oxidation of the VOCs in the 

3-36 
P.41 

catalyst bed will raise the temperature of the gases by an amount, AT, as 

the gases pass through the catalyst bed. An outlet temperature from the 

catalyst, and thus from the reactor, must be spedfied that will ensure the 

dedred level of destruction of the VOC stream. As in thermal indnerators, 

this value varies from compound to compound and also from catalyst to 

catalyst. Final design of the indnerator should be done by firms with experience in indnerator design. Guidelines given by Combustion Engineering 

[12] indicate that values from 300 to 900°F result in destruction efficiencies 

between 90 and 95 percent. To prevent deactivation of the catalyst a maximum bed temperature of 1,200°F should not be exceeded. In the example 

problem the catalyst outlet temperature, Tf., is sdected to be 900°F. 

Step 6c - Calculate the waste gas temperature at the exit of the 

preheater (primary) heat exchanger The waste gas temperature at 

the exit of the primary heat exchanger is estimated in the same manner as 

for the thermal indnerator. The equation for fractional energy recovery, 

Equation 3.16, is used, with the same assumptions as used for the thermal 

indnerator. For the example problem with a fractional energy recovery of 

0.70, a catalyst bed outlet temperature, T/., of 900°F, and a waste gas 

inlet temperature, TWi, of 100°F, the gas temperature at the exit of the 

preheater becomes 

r « . = 660°F 

The same considerations regarding the doseness of the temperature of 

the exhaust gas, T/#, to its dew point apply to the catalytic indnerator as 

they did to the thermal indnerator. 

Step 7c • Calculate the auxiliary fuel requirement, Qa/ The auxiliary fud requirement, QA/, is calculated by making mass and energy balances around the preheater combustion chamber and the catalyst chamber. 

The auxiliary fud requirement calculated in this manner must be checked to 

insure that it falls within the constraints imposed by design condderations 

of the catalytic indnerator. These constraints are as follows: 

• The auxiliary fud requirement must be podtive. A negative fuel 

requirement indicates that the heat of combustion of the waste gas, 

3-37 
P.42 

(—A/ie), is too high for the fractional energy recovery in the primary 

heat exchanger that was selected. 

e The auxiliary fud amount must be high enough to provide a stable 

flame in the preheater combustion chamber (See Step 8c below). 

An energy balance around the preheater combustion chamber and the 

catalyst chamber, taken together, results in Equation 3.18, the same equation used in the thermal indnerator calculations. The input data for Equation 3.18 for the catalytic indnerator example problem are summarized 

below: 

e The waste stream is essentially air so that 

Pv. - P»i = 0.0739 lb/scf, air at 77°F, 1 atm 

CPmair — 0.248 Btu/lb °F, the mean heat capadty of air between 77°F and 780°F(the average 

of the preheater exit and catalyst 

bed outlet temperatures) 

e Other input data to Equation 3.18 indude 

Qw. - Qwt= 20> 00°

 scf m 

(-AAe„) ss 21,502 Btu/lb, for methane 

Taf = 77°F, assume ambient conditions 

p9f = 0.0408 lb/ft3, methane at, 77°F, 1 atm 

Th s= 900°F, from Step 5c 

Tw. ss 660°F, from Step 6c 

(-Art,.) * 56.6 Btu/lb, from Step 4 

Substituting the above values into Equation 3.18 results in 

<?«/ ss 40 scfin 

If the outlet temperature of the catalyst bed, Tfn is 800° F, then Qa/ 

decreases to -6.7 scfm. In other words, no auxiliary fud would, theoretically, be required at this bed temperature. However, as discussed above in 

3-38 
P.43 

Step 8t, a certain quantity of auxiliary fuel would be required to maintain 

burner stability. 

At 70% energy recovery and 900°F outlet catalyst bed temperature, a 

waste gas with a heat of combustion, (-Arte,,), of about 79.9 Btu/lb would 

cause the auxiliary fud requirement, Qa /, to become negative, indicating 

the catalyst bed would exceed 900°F. At 70% energy recovery and 800°F 

outlet catalyst bed temperature, this same result occurs with a (-AACWS) of 

52.7 Btu/lb. Both of these heats of combustion are rdativdy low for typical 

waste gases. These results are, of course, dependent on the assumption of 

energy losses from the combustion chamber. The lower the energy losses, 

the lower the allowable waste gas heat of combustion before overheating 

occurs in the catalyst bed. 

Step 8c • Verify that the auxiliary fuel requirement is sufficient to 

stabilize the burner flame Only a small amount of auxiliary fud (< 5% 

of the total energy input) is needed to stabilize the burner flame. In general, 

more fuel than just this stabilizing fuel will be required to maintain the 

reactor temperature. It is wise to verify that the auxiliary fuel requirement 

calculated in Step 7c is suffident for stabilization. If it is insuffident, then a 

minimum amount of auxiliary fud must be used and the amount of energy 

recovery spedfied earlier must be reduced to avoid exceeding the specified 

temperature at which the indnerator will operate (Step 5c). 

This check is made in the same manner as that in Step 8t of the thermal 

indnerator calculation. The results of this check indicate that the auxiliary 

fud requirement is more than suffident to stabilize the burner flame. 

Step 9c • Estimate the inlet temperature to the catalyst bed, T,. 

The inlet temperature to the catalyst bed must be calculated to ensure 

that the inlet temperature is above that necessary to ignite the combustible 

organic compounds in the catalyst that was sdected for use. 

The inlet temperature to the catalyst bed, T,{, should be such that, 

when the temperature rise through the catalyst bed, AT, is added to it, 

the resulting temperature is Tfi, 900°F. Thus, 

AT = Tfi - Tu (3.21) 

3-39 
P.44 

The value of AT is determined by an energy balance around the preheater portion of the combustor. The preheater is required to heat the 

gases up to the catalyst bed inlet temperature udng auxiliary fuel.1 This 

energy balance is made with the assumptions made earlier in deriving Equation 3.18 and further assuming that only auxiliary fuel is combusted in the 

preheater portion. The resulting equation is very similar to Equation 3.18 

except that (1) the terms with an fi subscript become terms with rj subscripts to denote a catalytic reactor inlet stream rather than a combustor 

outlet (/lue gas inlet to the primary heat exchanger) and (2) the term 

for combustion of the waste gas organics does not appear. The resulting 

equation is as follows: 

*..Q«.[c>,.,..(i-ir,,. - TW. - o.ir,./)] 

p.tQ.t (-Art^-i.ic^o^-r,,,) = ...r;: ? z ; ;

f

 <3.») 

This equation may be rearranged to solve for Tr( explidtly. This produces an equation that is somewhat complex and non-intuitive. 

T p*tQ*A{-^K,) + i-icy^rTM/! + PV.Q-.C^.JT*. + o.ir,.,) 

l'lC9m^{pm,Qmt + p„.Qm.) 

(3.23) 

After substituting the example problem parameters into Equation 3.23, 

we obtain a value for Tr{ of 693°F. Based on ignition temperatures shown in 

Table 3.2, this reactor inlet temperature should be satisfactory. Prior to a 

more definitive design, the ignition temperatures for the spedfic chemicals 

should be verified. 

The temperature rise across the catalyst bed is thus (900 - 693) or 

207° F. These temperatures are somewhat send tive to the assumption for 

energy losses from the combustor. The assumption for energy losses is 

perhaps somewhat conservative, i.e., it causes a larger Q«/ to be estimated 

than would a less conservative assumption, and becomes more conservative 

as the combustor size and insulation are increased. 

1

 At equilibrium, the temperature of the catalyst bed is maintained without requiring 

auxiliary fuel. 

3-40 
P.45 

Step 10c - Calculate the total volumetric flow rate of gas through 

the incinerator, Q/{ The total volumetric flow rate of gas leaving the 

indnerator is referred to as the flue gas flow rate, Q/{, and is the gas rate on 

which the incinerator sizing and cost corrdations are based. The flue gas 

flow rate measured at the standard conditions of 77° F and 1 atmosphere, 

where the increase in volumetric throughput due to an increase in the number of moles of gas as a result of combustion is neglected, is the sum of the 

inlet streams to the indnerator. 

ss 20,000 + 0 + 40 

ss 20,040 scfm 

Step li e - Calculate the volume of catalyst in the catalyst bed 

If the volumetric flow rate of gas through the catalyst bed, Q/,, and the 

nominal residence time (redprocal space vdodty) in the catalyst bed are 

known, then the volume of catalyst can be estimated. There exists a complex set of relationships between the catalyst volume and geometry, overall 

pressure drop across the catalyst, converrion of the oxidizable components 

in the gas, gas temperature, and the reaction rate. These relationships are 

dependent on the catalyst and the type of compound bdng oxidized. It 

is beyond the scope if this Manual to discuss these rdationships, even in 

an approximate way. For the purposes of cost estimation, the space velodty, <p in redprocal time units, necessary to achieve the required level of 

destruction can be used to approximate the catalyst volume requirement. 

The space velodty is defined as 

where 

V^ s= Overall bulk volume of the catalyst bed, induding 

interpartide voids (ft3) 

By petro-chemical industry convention, the space vdodty is computed at 

the conditions of 60°F (not 77°F) and 1 atm. The volumetric flow rate, 

3-41 
Q/., must be corrected to these conditions. The proper space vdodty to 

achieve a desired levd of conversion is based on experimental data for the 

system involved. For predous metal monolithic catalysts, the space vdodty 

generally lies between 10,000 h_1 and 60,000 h"1. (Base metal catalysts 

operate at lower space velodties, ranging from 5,000 to 15,000 h~l.)[10] 

For the example, using a space vdodty of 30,000 h"1 or 500 min'1, and 

using QA at 60°F, 

/60 + 460N Qh at 608F = 20,040 ( l l "- 1 *" V77 + 460/ 

ss 19,400 ft3/min 

19,400 ft3/min 

"* ~ 500 min"1 

s= 39 ft3 

There are a number of catalyst bed parameters, such as catalyst configuration and bed design, that are not rignificant for study type cost estimates. 

Accordingly, design of these factors is not discussed here. 

• 

3.5 Cost Analysis 

This section presents procedures and data for estimating capital and annual 

costs for four types of indnerators:(l) thermal-recuperative, (2) thermalregenerative, (3) fixed-bed catalytic, and (4) fluid-bed catalytic. 

3.5.1 Estimatin g Tota l Capita l Investmen t , 

Total capital investment, TCI, indudes the equipment cost, EC, for the 

indnerator itself, the cost of auxiliary equipment {e.g., ductwork), all direct 

and indirect installation costs, and costs for buildings, dte preparation, offsite fadlities, land, and working capital. However, the last five costs usually 

do not apply to indnerators. (See Chapter 2 of this Manual for a detailed 

description of the dements comprising the TCI.) 

3-42 

P.46 

. 

u. 
P.47 

Table 3.7: Scope of Cost Correlations 

Total (Flue) Gas 

Indnerator Type Flowrate, scfm Figure Number 

Thermal - Recuperative 500a-50,000 3.4 

Thermal - Regenerative 10,000-100,000 3.5 

Fixed-Bed Catalytic 2,000-50,000 3.6 

Fluid-Bed Catalytic 2,000-25,000 3.7 

'Although Figure 3.4 coven the 1,000 to 50,000 scfm range, the correlation 

is valid for the SOO to 50,000 scfin range. 

3.5.1.1 Equipment Costs, EC 

As discussed in Section 3.2.3, the equipment costs, EC, given in this chapter apply to packaged indnerators, except for regenerative indnerators. 

For regenerative indnerators, the costs apply to fidd-erected units. The 

EC typically indudes all flange-to-flange equipment needed to oxidize the 

waste gas, induding the auxiliary burners, combustion chamber, catalyst, 

primary heat exchanger (except for the "zero heat recovery" cases), weathertight housing and insulation, fan, flow and temperature control systems, 

a short stack, and structural supports. Smaller units, e.g. typically less 

than 20,000 scfm, are typically preassembled and skid-mounted [18]. The 

various available indneration systems are presented in four groups delineated according to their similarity of design. These groups are outlined in 

Table 3.7. With the exception of regenerative thermal and fluid-bed catalytic indnerators, the maximum size for which costs are given is 50,000 

scfm. Although larger units of each technology can be built, applications 

are rare at flow rates above 50,000 scfm. Regenerative thermal incinerator 

costs are provided for flow rates from 10,000 to 100,000 scfm. Fluid-bed 

catalytic indnerator costs are provided for flow rates from 2,000 to 25,000 

scfm. 

The cost curves are least-squares regressions of cost data provided by 

different vendors. It must be kept in mind that even for a given indneration technology, design and manufacturing procedures vary from vendor 

to vendor, so that costs may vary. As always, once the study estimate is 

J

 3"43 
P.48 

completed, it is recommended that more than one vendor be solidted for a 

more detailed cost estimate. 

The additional expense of add gas clean-up or particulate control is not 

treated in this section. The equipment cost of a gas absorber to remove any 

add gases formed in the indnerator can be quite large, sometimes exceeding 

the equipment cost of the indnerator itself even for simple packed tower 

scrubbers [19]. For more complex absorbers that indude venturi scrubbers 

instead of, or in addition to, packed beds, the cost of the scrubber alone 

may be up to 4 times that of the incinerator [11]. These more complex 

absorbers are sometimes necessary when particulates, in addition to add 

gases, must be removed from the flue gas. 

Thermal Incinerators Among the thermal units, the direct flame (0% 

energy recovery) and recuperative systems are treated together because 

the various levels of energy recovery are achieved simply by adding heat 

exchanger surface area. Costs for these units were provided by several vendors [12,20,21]. The EC of these units are given as a function of total volumetric throughput, Qt*, in scfm. "Qtot" is the total volume of the gaseous 

compounds exiting the combustion chamber; it is identical to the term, 

uQ/{", used in Figures 3.1 and 3.3. This indudes the combustion products, 

nitrogen, unbumed fud and organics, and other constituents. (See Figure 

3.4). Note that costs are given free on board (F.O.B.) in April 1988 dollars. 

Based on a least-squares regression analyris, a log-log rdationship between 

throughput and EC was found for a given levd of energy recovery (HR) 

over the flow rate range from 500 to 50,000 scfm. These relationships are 

as follows: 

EC = 10294Q£f" HR = 0% (3.24) 

EC =s 13149QJ2609 HR = 35% (3.25) 

EC = 17056(?^SM HR = 50% (3.26) 

EC = 21342g£JMO HR=70% (3.27) 

The regenerative (or excess enthalpy) systems provide up to 95 percent 

heat recovery at the expense of higher capital costs. Thdr unique design 

[22,23], which combines the heat exchanger and reactor, is substantially 

different from traditional thermal units and is therefore treated separately 

in Figure 3.5. The ECs of these systems are given as an approximatdy 

3-44 
P.49 

Q 

< 

in 

o X 

1— 

in 

cc 

900 

800 

700 

600 

SOO 

400 

300 

o a 

oo 

00 

cn 

a: 

a. 

< 

aa" 

o 

to 

O 

O 

LU 

0. 

D 

a 

Ld 

200 

100 

90 

80 

70 

so 

50 

40 

30 

20 

10 

1 

* } 

i 

! 

1 

i 

i 

. 0 0 X ENERGY RECOVERY 

•DM! ! 

. A 50 X 

• 70 X 

i 

4 S S 7 8 9 10 20 30 40 SO 60 708090 

FLUE GAS VOLUMETRIC FLOW RATE, SCFM (THOUSANDS) 

Figure 3.4: Equipment Costs of Thermal Indnerators, Recuperative 

3-45 
P.50 

v> 

Q 

<fl 3 O 

jE 

tfl 

o 

Q 

00 

OO 

0 1 

03 

O 

U. 

(fl 

o 

o 

UJ 

2 

0 . 

O 

Ui 

o 

o 

C— N 

o 

CO 400 

o 

0 

o 95 % ENERGY RECOVERY 

»• .0 i r0 6 0 

. i 

-:•• 

80 100 

FLUE GAS VOLUMETRIC FLOW RATE, SCFM (THOUSANDS) 

Figure 3.5: Equipment Costs of Thermal Indnerators, Regenerative 

3-46 
P.51 

linear function of total flow rate over a 10,000 to 100,000 scfm range by the 

following equation: 

EC = 2.204 x IO8 + 11.57 (?«* (3.28) 

Again, the higher capital costs of these units can be substantially offset 

by the substantial savings in auxiliary fud costs. 

Catalytic Incinerators The EC for a catalytic indnerator is a function 

of the type of catalyst contacting pattern used and the total gas flow rate, 

Qtot, for a given level of energy recovery. There are three types of contacting 

configurations used in catalytic systems: fixed-bed, catalytic monolith, and 

fluid-bed. The EC for the first two are generally comparable and are given 

in Figure 3.6. The data provided by several vendors [12,20,21,24] exhibited 

curvilinear rdationships with Qu* for each of the energy recovery rates. 

Least squares regressions of the data yidded the following correlations for 

total flow rates between 2,000 and 50,000 scfm: 

EC 

EC 

EC 

EC 

= 1105QTM4n 

= 3623<?£41W 

= 1215C?^8T8 

s= 1443g£8a7 

HR =s 0 % 

HR = 35 % 

HR = 50 % 

HR = 70 % 

(3.29) 

(3.30) 

(3.31) 

(3.32) 

Fluid-bed catalytic indnerators afford certain advantages over fixed-bed 

catalyst units in that they tolerate waste streams with (1) higher heating 

values, (2) particulate contents, and (3) chlorinated spedes. For this enhanced flexibility of feed streams, a higher capital cost is incurred, as indicated by the EC shown in Figure 3.7. The data shown were provided 

by vendors [11,19] and exhibited a linear rdationship over the range of 

flow rates from 2,000 to 25,000 scfm. They can be approximated by the 

following equations: 

EC = 8.48 x 104 + 13.2C?** HR = 0 % (3.33) 

EC = 8.84 x 104 + U.6QM HR = 35 % (3.34) 

EC = 8.66 x 104 + 15.8*?** HR = 50 % (3.35) 

EC - 8.39 x 104 + 19.2C?*, HR = 70 % (3.36) 

A comparison of the EC of thermal, catalytic fixed-bed, and catalytic 

fluid-bed systems with 50 percent energy recovery is shown in Figure 3.8. 

3-47 
P.52 

tn 

o 

z 

< 

in 

O 

X 

1— 

to 

OL 

5 _ j 

o 

o 

GO 

00 

cf M n 

_l 

ca CL 

< 

CD 

O 

u_ 

>-

o 

CJ 

h -

z UJ 

2 

a. 

D 

O 

900 

800 

700 

600 

500 

400 

300 

200 

100 

90 

80 

70 

60 

50 

40 

30 

20 

10 

i 

-

'V^ 

• ^ 

?x* V: 

I 1 

I 1 

1 | 

••••• ; .-" . - " 1 - • 

• = • • ' y^ ^~^ 

-

0 0 % ENERGY RECOVERY 

a 35 % 

A 50 % " " 

• 70 X 

• 

4 5 6 7 8 9 10 20 30 40 50 60 70 8090 

FLUE GAS VOLUMETRIC FLOW R E, SCFM (THOUSANDS) 

Figure 3.6: Equipment Costs of Catalytic Indnerators, Fixed-Bed 

3-48 
Y-Ht^^Bv^ v ^^FV^ ^ 

P.53 

(fl 

a 

o 

X 

tfl 

o 

a 

ao 

ao 

OS 

k 

O 

Lu 

(fl 

o o. 

z* 

Ul 

2 

QL 

O 

Ul 

600 

o 

o 

^* 

o 

CN 

o 

o 0 % ENERGY RECOVERY 

• 35 % " 

A 50 % " 

• 70 % " 

• 

10 20 30 40 50 

FLUE GAS VOLUMETRIC. FLOW RATE. SCFM (THOUSANDS) 

Figure 3.7: Equipment Costs of Catalytic Incinerators, Fluid-Bed 

3-49 
P.54 

cn 

a z 

< 

cn 

o X 

Hcn 

a: 

5 

_ i 

o 

a 

03 

ao 

cn ^» 

• _ i 

cr 

a. < 

CD 

O 

u. 

Hcn 

o 

o 

z 

UJ 

2 

CL 

3 

O 

900 

800 

700 

600 

500 

400 

300 

200 

100 

90 

80 

70 

60 

50 

40 

30 

20 

i n 

i 

1 

1 —- i 

1 1 i 

! 

1 

i 

i 

. 0 CATAL 

. A THERI 

50% Eh 

* • r 

1 

YUC - FLUI0 BED 

YT1C - FIXED BED 

JAL - RECUP. 

^ERGY RECOVERY 

! 

4 5 6 7 8 9 10 20 30 40 SO 60 708090 

FLUE GAS VOLUMETRIC FLOW RATE. SCFM (THOUSANDS) 

Figure 3.8: Equipment Cost Comparison of Indnerator Types 

3-50 
P.55 

-TIP "W^TP •^r,"~—rr~ - ^ ' 

3.5.1.2 Installation Costs 

As explained in Chapter 2, the purchased equipment cost, PEC, is calculated by taking the sum of the EC and the cost of auxiliary equipment 

{e.g., ductwork), taxes, freight, and instrumentation. Average values of 

direct and indirect installation factors [25] to be applied to the PEC are 

given in Table 3.8 for both recuperative thermal and fixed- and fluid-bed 

catalytic indnerators. 

Table 3.9 shows the itemized installation costs that are obtained when 

these installation factors are applied to the PECs for the example indnerators. Depending on the site conditions, the installation costs for a given 

indnerator could deviate significantly from costs generated by these average factors. Vatavuk and Neveril [25] provide some guidelines for adjusting 

the average installation factors to account for other-than-average installation conditions. For units handling total gas flow rates lower than 20,000 

scfm the installation costs are. minimal, amounting normally to only utility tie-ins (electrical and, if necessary, combustion or dilution air). The 

installation costs for these smaller indnerators would be 20 to 25 % of the 

PEC. Smaller units may be installed on the roofs of manufacturing buildings rather than at ground levd. In such cases the installation factors could 

be as high as (or higher than) the factors shown in Table 3.8, even though 

the units would be "packaged". 

3.5.2 Estimating Total Annual Cost 

The total annual cost (TAC) is the sum of the direct and indirect annual 

costs. The TAC for both example systems is given in Table 3.10, along 

with suggested factors for calculating them. 

3.5.2.1 Direct Annual Costs 

Direct annual costs for indnerators indude labor (operating and supervisory), maintenance (labor and materials), fud, electridty, and (in catalytic 

units) replacement catalyst. For thermal and catalytic units, the fud usage 

rate is calculated as shown in Sections 3.4.2 and 3.4.3, respectively, where 

3-51 
P.56 

Table 3.8: Capital Cost Factors for Thermal and Catalytic Indnerators" 

Cost Item 

Direct Costs 

Purchased equipment costs 

Indnerator (EC) + auxiliary equipment* 

Instrumentation0 

Sales taxes 

Frdght 

Purchased equipment cost, PEC 

Direct installation costs 

Foundations & supports 

Handling & erection 

Electrical 

Piping 

Insulation for ductwork'' 

Painting 

Direct installation cost 

Site preparation 

Buildings 

Total Direct Cost, DC 

Indirect Costs (installation) 

Engineering 

Construction and Add expenses 

Contractor fees 

Start-up 

Performance test 

Contingendes 

Total Indirect Cost, IC 

Total Capital Investment = DC + IC 

Factor 

As estimated, A 

0.10 A 

0.03 A 

0.05 A 

B a 1.18 A 

0.08 B 

0.14 B 

0.04 B 

0.02 B 

0.01 B 

0.01 B 

0.30 B 

As required, SP 

As required, Bldg. 

1.30 B + SP + Bldg. 

0.10 B 

0.05 B 

0.10 B 

0.02 B 

0.01 B 

0.03 B 

0.31 B 

1.61 B + SP + Bldg. 

* Reference (38j. 

'Ductwork tad aay other equipment normally not included with unit famished by indnerator vendor. 

'Instrumentation aad control* often famished with the incinerator, aad Una often included ia tha BC. 

4

If ductwork dimensions have beea estabUshed, coat msy be estimated baaed oa tlO to S13/ft* of 

surface for field application. Fan housings cad stock* mop cloo bo iaselated. 

3-52 
P.57 

Table 3.9: Capital Costs for Thermal and Catalytic Indnerators 

Example Problem 

Cost, S 

Cost Item Thermal- Fluid-Bed 

Recuperative Catalytic 

Direct Coats 

Purchased equipment costs 

Incinerator (EC) 1254,200 $468,200 

Auxiliary equipment4 

Sum = A 

Instrumentation, 0.1 A 

Sales taxes, 0.03A 

Freight, 0.05A 

Purchased equipment cost, B 

Direct installation costs 

Foundation and supports, 0.08B 

Handling and erection, 0.14B 

Electrical, 0.04B 

Piping, 0.02B 

Insulation (for ductwork), 0.01B 

Painting, 0.01B 

Direct installation cost $90,000 $165,600 

Site preparation4 

Buildings4 

$254,200 

25,400 

7,630 

12,700 

$300,000 

24,000 

42,000 

12,000 

6,000 

3,000 

3,000 

$468,200 

46,800 

14,000 

23,400 

$552,400 

44,200 

77,300 

22,100 

11,000 

5,520 

5,520 

Total Direct Cost 

Indirect Costs (installation) 

Engineering, 0.10B 

Construction-and field expenses, 0.05B 

Contractor fees, 0.10B 

Start-ap, 0.02B 

Performance test, 0.01B 

Contingencies, 0.03B 

Total Indirect Cost 

Total Capital Investment (rounded) 

$390,000 

30,000 

15,000 

30,000 

6,000 

3,000 

9,000 

$93,000 

$483,000 

$718,000 

55,200 

27,600 

55,200 

11,000 

5,520 

16,600 

$171,100 

$889,000 

.'None of these items ia required. 

3-53 
P.58 

Table 3.10: Annual Costs for Thermal and Catalytic Indnerators 

Example Problem 

Cost Item 

Direct Annual Costa*,DC 

Operating Labor 

Operator 

Supervisor 

Operating materials 

Maintenance 

Labor 

Material 

Catalyst replacement 

Utilities 

Natural Gas 

Electricity 

Total DC 

Indirect Annual Costs, IC 

Overhead 

Administrative charges 

Property taxes 

Insurance 

Capital recovery* 

Suggested Factor 

0.5 h/shift 

15% of operator 

—. 

0.5 h/shift 

100% of 

maint. labor 

100% of catalyst 

replaced ea 2 yr 

— 

— 

60% of sum of 

operating, supv., 

k maint. labor Jc 

maint. materials. 

2% TCI 

1% TCI 

1%TCI 

CRF [TCI -

1.08 (Cat. Cost)] 

Unit Cost* 

$12.96/h 

— 

$14.26/h 

^~ 

3650/ft3 for 

metal oxide 

$3.30/kft* 

$0.059/kWh 

-

— 

— 

— 

-snesa 

Thermal 

6,480 

973 

7,130 

7,130 

0 

264,500 

35,000 

$321,200 

13,000 

9,660 

4,830 

4,830 

78,600 

Fluid-Bed 

Catalytic 

6,480 

* 972 

7,130 

7,130 

14,600 

83,400 

42,300 

3142,000 

13,000 

i. 17,800 

- 8,900 

,„ 8,900 

"142,200 

Total IC 8110,900 8190,800 

Total Annual Coat (rounded) 8432,000 $352,000 

*19SS dollars. 

'Assumes 8,000 h/yr. 

'The capital recovery cost factor, CRF, ia a (unction of the catalyst or equipment life (typically, 3 and 10 

years, respectively) and the opportunity coot of the capital («.«., interest rate). For example, for a 10 year 

equipment life and a 10% interest rate, CRF s 0.1838. 

3-54 
P.59 

Table 3.11: Typical Pressure Drop Across Selected Equipment 

Equipment Type 

Thermal Indnerators 

Catalytic Fixed-bed Indnerators 

Catalytic Fluid-bed Indnerators 

Heat exchangers 

n n 

» » 

Energy Recovery, % 

0 

0 

0 

35 

50 

70 

A P , in. H,0 

4 

6 

6-10 

4 

8 

15 

natural gas (methane) is assumed to be the fud. (Other fuels could be used 

for thermal units.) 

The electridty costs are primarily assodated with the fan needed to 

move the gas through the indnerator. The power (in kilowatts) needed 

to move a given volumetric flow rate of air {Qu* per Sections 3.4.2 and 

3.4.3) at a total flange-to-flange pressure drop of A P inches of water and 

combined motor/fan effidency, e, is adapted from Equation 2-7, as follows: 

Paw„lm _ 1-1? * 10-<Q,.,AJ> ( J 3? ) 

Fan effirienries vary from 40 to 70 percent [15] while motor effidendes are 

usually 90 percent. 

The total pressure drop across an indnerator system depends on the 

number and types of equipment elements induded in the system and on 

design condderations. The estimation of actual pressure drop requirements 

involves complex calculations based on the spedfic system's waste gas and 

flue gas conditions and equipment used. For the purposes of this section, 

however, the approximate values shown in Table 3.11 can be used. 

For the example cases, we will assume 8,000 hours per year operation 

and a 60% effidency for the fan and motor together. Udng pressure drops of 

4 and 8 inches of water, respectively, for the thermal and fluid-bed catalytic 

indnerators3, and adding the pressure drop of 15 inches of water for 70% 

'A fluid-bed catalytic incinerator is used because the waste gas contains a chlorinated 

compound which would poison the catalyst in a fixed-bed incinerator. 

3-55 
P.60 

heat recovery, the fan power requirements can be calculated as follows: 

Thermal Indnerator 

_ 1.17 x 10~4(20,000 sc£m)(19 inches water) 

Power"» = 61b~ 

= 74.1 kW 

Catalytic Indnerator 

_ 1.17 x 10"4(20,000 scfm)(23 inches water) 

P9W€r^ =

 — b^o 

= 89.7 kW 

The annual dectridty costs would be the products of these usages, the 

annual operating hours, and the dectridty cost (S/kWh), or: 

Electridty Cost (Thermal) = 74.1 kW x 8,000 hours/yr x $0.059/kWh 

= .35,000 per yr 

Electridty Cost (Catalytic) s 89.7 kW x 8,000 hours/yr x $0.059/kWh 

= 142,300 per yr 

V 

The catalyst replacement costs and scheduling are highly variable and 

depend on the nature of the catalyst, the amount of "poisons" and particulates in the gas stream (including the auxiliary fud), the temperature 

history of the catalyst, and the design of the unit. It is impossible to predict the costs in a general sense. However, noble metal monolith catalysts 

operating on pure hydrocarbon gases in air will last longer than fluid-bed 

base metal catalysts operating on chlorinated hydrocarbons in air. Noble 

metal catalysts are also more expensive than base metal oxide catalysts. 

The catalyst life for many fidd units is from 1 to 4 years. The cost, in 

April 1988 dollars, of the replacement catalyst must be obtained from the 

vendor, but it may be estimated at $3,000/ft3 for noble metal catalysts and 

$650/ft3 for base metal oxide catalysts. For the example case, the catalyst 

is a base metal oxide because the waste gas contains a chlorinated compound. We will assume a two year catalyst life. Knowing that the catalyst 

volume is 39 ft3 (Section 3.4.3) and using a cost of $650/ft3 and a capital 

3-56 
P.61 

recovery factor of 0.5762 (2-year life at a 10% interest rate), the annual 

expense for catalyst replacement is 

Annual Catalyst Replacement Cost = 39 ft3 x 650 p x 0.5762 

= $14,600 per year 

To calculate the fuel or dectridty annual cost, multiply the fuel usage 

rate (scfm) or the electridty usage rate (kW) by the total hours per year of 

operation {e.g., 333 d/yr x 24 h/d = 8,000 h/yr) and by the appropriate 

unit cost {e.g., S/sc£m for fud and S/kWh for dectridty). 

For the example cases, the fud costs can be calculated from the fuel 

usage rates and the natural gas unit cost of S0.00330 /scf.. For the thermal 

indnerator example, the annual fud cost is calculated as follows: 

Annual Fud - «*— $ .«.- scf „,,min „ ^^hr 

.T- , . = 0.00330— x 167-r- x 60-r-- x 8,000— 

v^ost, lnermal scf mm hr yr 

= $264,500 per yr 

For the catalytic indnerator example, the annual fud cost is found 

similarly: 

Annual Fud Cost, Catalytic = $63,400 per year 

Operating and maintenance labor are estimated as 0.5 hours per 8-hour 

shift each, supervisory labor at 15 % of operating labor, and maintenance 

material as 100 % of maintenance labor. 

3.5.2.2 Indirect Annual Costs 

The indirect (fixed) annual costs indude capital recovery, overhead, and 

property taxes, insurance, and administrative (G&A) charges. The last 

three of these can be estimated at 1%, 1%, and 2% of the total capital 

investment, respectivdy. The system capital recovery cost is based on an 

estimated 10-year equipment life. (See Section 2 for a thorough discussion 

of the capital recovery cost and the variables that determine it.) The system 

3-57 
P.62 

capital recovery cost is the product of the system capital recovery factor 

(CRF) and the total capital investment (TCI) less the purchased cost of 

the catalyst (Cost x 1.08 where the 1.08 is for frdght and sales tax). These 

values calculated for the example cases are given in Table 3.10. 

3.5.3 Cost Comparison for Example Case 

The example v OC stream defined in Section 3.4.1 serves to illustrate some 

typical characteristics of thermal and catalytic systems. The total annual 

costs shown in Table 3.10 show that the catalytic system's auxiliary fud 

costs are significantly lower than those of the thermal unit. The disparity 

is enough to offset the higher capital costs of the catalytic indnerator over 

the assumed 10-year lifetime of the units. Two factors that should be 

noted in the comparison of these two systems are (1) the 98 percent levd 

of destruction met by the thermal indnerator may be difficult to reach by 

the catalytic system (this may be important in some cases), and (2) the 

example waste stream is of particularly low heating value (4 Btu/scf) which 

favors the catalytic system due to the lower auxiliary fud requirements. 

3.6 Acknowledgements 

The authors gratefully acknowledge the following companies for contributing data to this chapter: 

• Peabody Engineering (Stamford, CT) 

• Combustion Engineering - Air Preheater, Inc. (Wellsville, NY) 

• TEC Systems, Inc (DePere, WI) 

• Air Research, Inc. (ARI) (Palatine, IL) 

• Energy Development Assodates (EDA) (Itasca, IL) 

• Pillar Technologies, Inc. (Hartland, WI) 

• Huntington Energy Systems, Is c (Unioe, NJ) 

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P.63 

• Regenerative Environmental Equipment Co. (REECO) (Morris Plains, 

NJ) 

• Englehard Corp. (Edison, NJ) 

3-59 
P.64 

Appendix 3A 

Properties of Selected 

Compounds 

% 3-60 
P.65 

Table 3.12: Limits of Flammability of Combustible Organic Compounds 

in Air at Atmospheric Pressure, Room Temperature* 

Compound 

Methane 

Ethane 

Propane 

Butane 

Pentane 

Hexane 

Octane 

Nonane 

Decane 

Ethylene 

Propylene 

Acetylene 

Cydohexane 

Benzene* 

Toluene 

Molecular 

Wdght 

16.04 

30.07 

44.09 

58.12 

72.15 

86.17 

114.23 

128.25 

142.28 

28.05 

42.08 

26.04 

84.16 

78.11 

92.13 

LELa, 

vol. % 

5.00 

3.00 

2.12 

1.86 

1.40 

1.18 

0.95 

0.83 

0.77 

2.75 

2.00 

2.50 

1.26 

1.40 

1.27 

UEL*, 

vol. % 

15.00 

12.50 

9.35 

8.41 

7.80 

7.40 

28.60 

11.10 

80.00 

7.75 

7.10 

6.75 

'Reference [14] 

"Lower Explosive Limit 

*Upper Explosive Limit 

3-61 
P.66 

Table 3.13: Molar Heat Capadties of Gases at Zero Pressure * 

C ^ a + aT + cT' + dT3; Tin'K 

fl C.aT 

"• (Ta - TO 

Cp in calories/g-mole °K or Btu/lb-mole °R 

\ 

Compound 

Methane 

Ethane 

Propane 

Butane 

Pentane 

Hexane 

Cyclopentane 

Cydohexane 

Benzene 

Toluene 

Nitrogen 

Oxygen 

Air 

Carbon dioxide 

a 

4.750 

1.648 

-0.966 

0.945 

1.618 

1.657 

-12.957 

-15.935 

-8.650 

-8.213 

6.903 

6.085 

6.713 

5.316 

b xl0a 

1.200 

4.124 

7.279 

8.873 

10.85 

13.19 

13.087 

16.454 

11.578 

13.357 

-0.03753 

0.3631 

0.04697 

1.4285 

c xlO* 

0.3030 

-1.530 

-3.755 

-4.380 

-5.365 

-6.844 

-7.447 

-9.203 

-7.540 

-8.230 

0.1930 

-0.1709 

0.1147 

-0.8362 

d xlO9 

-2.630 

1.740 

7.580 

8.360 

10.10 

13.78 

16.41 

19.27 

18.54 

19.20 

-0.6881 

0.3133 

-0.4696 

1.784 

Temperature 

Range,

 9K 

273-1500 

273-1500 

273-1500 

273-1500 

273-1500 

273-1500 

273-1500 

273-1500 

273-1500 

273-1500 

273-1800 

273-1800 

273-1800 

273-1800 

'Reference (26) 

3-62 
P.67 

Table 3.14: Heats of Combustion of Sdected Gaseous Organic Compounds, —A/ie, at 25°C and constant pressure to form gaseous water and 

carbon dioxide.' 

Compound 

Methane 

Ethane 

Propane 

Butane 

Pentane 

Hexane 

Octane 

Nonane 

Decane 

Ethylene 

Propylene 

Cydopentane 

Cydohexane 

Benzene 

Toluene 

Molecular 

Weight 

16.04 

30.07 

44.09 

58.12 

72.15 

86.17 

114.23 

128.25 

142.28 

28.05 

42.08 

70.13 

84.16 

78.11 

92.13 

-AA, 

cal/g. 

11,953.6 

11,349.6 

11,079.2 

10,932.3 

10,839.7 

10,780.0 

10,737.2 

10,680.0 

10,659.7 

11,271.7 

10,942.3 

10,563.1 

10,478.7 

9,698.4 

9,784.7 

Btu/lb 

21,502 

20,416 

19,929 

19,665 

19,499 

19,391 

19,256 

19,211 

19,175 

20,276 

19,683 

19,001 

18,846 

17,446 

17,601 

•Reference [IS] 

3-63 
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3-64 
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3-65 
P.70 

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