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OAQPS CONTROL COST MANUAL
Fourth Edition
EPA 450/3-90-006
January 1990
United States Environmental Protection Agency
Office of Air Quality Planning and Standards
Research Triangle Park, North Carolina 27711
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1. REPORT NO.
EPA 450/3-90-006b
I. R
E
f
rjOA
TECHNICAL REPORT DATA
(flease read Instructions on the reverse before completing)
1. TITLE AND SUBTITLE
2.
OAQPS Control Cost Manual (Fourth Edition): Supplement 2
3. RECIPIENT'S ACCESSION NO.
5. REPORT DATE
'October 1992
. PERFORMING ORGANIZATION CODE
Radian: W. Barbour, R. OOmmen, G.Shareef
EPA: W. Vatavuk
8. PERFORMING ORGANIZATION REPORT NO.
9. PERFORMING ORGANIZATION NAME AND ADDRESS'
Radian Corporation
P.O. Box 13000
Research Triangle Park, NC 27709
10. PROGRAM ELEMENT NO.
11. CONTRACT/GRANT NO.
EPA-68-D1-0117 (W. A. 20)
12. SPONSORING AGENCY NAME AND ADDRESS
Environmental Protection Agency
Office of Air and Radiation
Office of Air Quality Planning and Standards
Research Triangle Park, NC 27711
3. TYPE OF REPORT AND PERIOD COVERED
Final
4. SPONSORING AGENCY CODE
This is the second supplement to the OAQPS Control Cost
Manual (Fourth Edition). The supplement consists of a new Manual
chapter, Chapter 9 ("Gas Absorbers"). Like the other chapters in
the Manual, Chapter 9 is self-contained. It discusses: (1) the
types and applications of packed column gas absorbers used in air
pollution control; (2) the theory underlying their operation and
design; (3) basic sizing procedures; and (4) current data and
procedures for estimating study-level (+ 30%-accurate) capital
and annual costs, in particular, the chapter contains 1991
column and packing costs, which are correlated with appropriate
sizing parameters (e.g., column height and diameter). Finally,
Chapter 9 includes: a comprehensive example problem that
illustrates the sizing and costing procedures; three appendices;
a table of contents; and a list of references.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
Stationary emission sources
Costs
Control techniques
Control device design/sizing
Gas absorbers
Packed columns
Packing
". DISTRIBUTION STATEMENT
Unlimited
b.lDENTIFIERS/OPEN ENDED TERMS
Cost estimation
Capital costs
Equipment, installation
Annual costs (direct,
Operating and maintenance
"Add-on" controls
19. SECURITY CLASS (Tilts Report)
Unclassified
20. SECURITY CLASS (This page I
Unclassified
c. COSATI Field/Group
casts
indirect)
costs
21. NO. OF PAGES
67
22. PRICE
I
Form 2220-1 (Rev. 4-77) PREVIOUS EDITION is OBSOLETE
I1 S. r';.v:^n;rionta! Protection Agency
.;;:ck:;on Bcu:?v-,,'-!, 12th Floor
iL 60604-3690
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INSTRUCTIONS
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Include a brief (200 \vords or less) factual summary of the most significant information contained in the report. It Hit- report ioniums a
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17. KEY WORDS AND DOCUMENT ANALYSIS
(a) DESCRIPTORS - Select from the Thesaurus of Engineering and Scientific Terms the proper autluiri/ed li-rms llul identity the major
concept of the research and are sufficiently specific and precise to be used as index entries I'or cataloging.
(b) IDENTIFIERS AND OPEN-ENDED TERMS - Use identifiers for project names, code names, equipment designjtors, etc. Use open-
ended terms written in descriptor form for those subjects for which no descriptor exists.
(c) COSATI I H'-LD GROUP - Held and group assignments are to be taken from the 1965 C'OSAl 1 Subject ( atepory List. Since the ma-
jority of documents are multidisciplmary in nature, the Primary l-ield/Group assignment^) will be spculii. discipline, jrea ol human
endeavor, or type of physical object. The application^) will be cross-referenced with sciondary 1 ickl/(,roiip assignments (lul will lollou
the primary postmg(s).
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Denote releasability to the public or limitation for reasons other than security (or example "Release Ijnliiinleil." ( He JMV avjil.ihiltiy In
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21. NUMBER OF PAGES
Insert the total number of pages, including this one and unnumbered pages, but exclude distribution list, il any.
22. PRICE
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EPA Form 2220-1 (Rev. 4-77) (Reverse)
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UNITED STATES ENVIRONMENTAL PROTECTION AGENCY
Office of Air Quality Planning and Standards
Research Triangle Park, North Carolina
Dear Manual Requester:
f-f'L '•/:;/". 3 7''1 ''"'"'<• C,
Enclosed is a copy of the third supplement to the OAQPS
Control Cost Manual (Fourth Edition). This supplement consists
of a new chapter, Chapter 10 ("Hoods, Ductwork, and Stacks").
Like the parent report, Supplement 3 is unbound to make it easier
for you to insert the pages into the same 3-ring binder that
contains your copy of the Manual.
Like the other Manual chapters, Chapter 10 is
self-contained. It discusses: (1) the types of hoods, ductwork,
and stacks used to support add-on air pollution control devices,-
(2) the theory underlying their operation and design; (3) basic
sizing procedures; (4) procedures and current data for estimating
capital and annual costs; and (5) several example problems that
illustrate the various sizing and costing procedures. Chapter 10
also contains a table of contents and a list of references.
To request additional copies of this supplement or any other
parts of the Manual, please telephone the Emission Standards
Division Control Technology Center (CTC) at (919) 541-0800. You
have already been placed on the mailing list and will receive
forthcoming Manual supplements as they are completed.
If you have any questions about the Manual, please telephone
me at (919) 541-5309 or telefax me at (919) 541-4028.
Thank you for your interest in the Manual.
Yours truly,
William M. Vatavuk, P.E.
Cost and Economic Impact Section
Standards Development Branch
Emission Standards Division
Enclosure
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Abstract (Item #16) for EPA Form 2220-1
This is the third supplement to the OAQPS Control Cost .
Manual (Fourth Edition). The supplement consists of a new Manual
chapter, Chapter 10 ("Hoods, Ductwork, and Stacks") Like the
other chapters in the Manual, Chapter 10 is self-contained. It
Hi trusses- (D the types and applications of hoods used to
support add-on air pollution control devices; (2) the theory
underlying their operation and design; (3) basic sizing
• and (4) procedures and current (1993) data for
qate
the prices of each type of equipment reixeuu cxu .L <=«»._ ..»« —~~
of fabrication materials, such as carbon and 304 stainless steel
?nlate and sheet types), FRP (fiberglass-reinforced plastic), and
P?C^poiyvinyl chloride . These prices have been correlated with
appropriate sizing parameters (e.g., duct ^terK Finally -
chanter 10 includes several example problems that illustrate tne
varies sizing and costing procedures; a table of contents; and a
list of references.
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TECHNICAL REPORT DATA
(Please read Instructions on the reverse before completing)
1. REPORT NO.
EPA 450/3-90-006c
3 RECIPIENT'S ACCESSION NO
TITLE AND SUBTITLE
JAQPS Control Cost Manual (Fourth Edition):Supplement J
5. REPORT DATE
March 1994
6. PERFORMING ORGANIZATION CODE
7. AUTHOR(S)
William M. Vatavuk
8. PER
iT NO.
9. PERFORMING ORGANIZATION NAME AND ADDRESS
Environmental Protection Agency
Office of Air and Radiation
Office of Air Quality Planning and Standards
Research Triangle Park, NC 27711
10 PROGRAM ELE
11 CONTRACT/GRANT NO.
12. SPONSORING AGENCY NAME AND ADDRESS
Environmental Protection Agency
Office of Air and Radiation
Office of Air Quality Planning and Standards
Research Triangle Park, NC 27711
13. TYPE OF REPORT AND PERIOD CO'
Final
14. SPONSORING AGENCY CODE
15. SUPPLEMENTARY NOTES
16. ABSTRACT
SEE ATTACHED
17.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b IDENTIFIERS/OPEN ENDED TERMS
Stationary emission sources
Costs
Control techniques
Control device design/sizing
Auxiliary equipment
Hoods
Ductwork
Stacks
. DISTRIBUTION STATEMEN"
Unlimited
~EPA Form 2220-1 (R«v. 4-77) PREVIOUS EDI T,ON i s OBSOLE -
Cost estimation
Capital costs
Equipment, installation Co
Annual costs (direct,indir
Operating and miantenance
"Add-on" controls
its
ect)
costs
19 SECURITY CLASS
Unclassified
s Reporri
20 SECURITY CLASS (This paij
Unclassified
COSATl field/Group
21 NO OF PAGES
6J
22 PRICE"
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This fourth edition of the OA QPS Control Coat 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.
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Chapter 1
INTRODUCTION
William M. Vatavuk
Standards Development Branch, OAQPS
U. S. Environmental Agency
Research Triangle Park, NC 27711
November 1989
Contents
1.1 Role of Cost in Setting of Regulations 1-2
1.2 Purpose of Manual 1-2
1.3 Organization of the Manual 1-3
1.4 Intended Users of the Manual 1-4
1.5 "Uniqueness" of the Manual 1-5
References 1-7
1-1
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1.1 Role of Cost in Setting of Regulations
Cost has an important role in setting many state and federal air pollution
control regulations. The extent of this role varies with the type of regula-
tion. For some types of regulations, cost is explicitly used in determining
their stringency. This use may involve a balancing of costs and environ-
mental impacts, costs and dollar valuation of benefits, or environmental
impacts and economic consequences of control costs.
For other types of regulations cost analysis is used to choose among
alternative regulations with the same level of stringency. For these regula-
tions, the environmental goal is determined by some set of criteria which
do not include costs. However, cost-effectiveness analysis is employed to
determine the minimum cost way of achieving the goal.
For some regulations, cost influences enforcement procedures or require-
ments for demonstration of progress towards compliance with an air quality
standard. For example, the size of any monetary penalty assessed for non-
compliance as part of an enforcement action needs to be set with awareness
of the magnitude of the control costs being postponed by the noncomplying
facility. For regulations without a fixed compliance schedule, demonstra-
tion of reasonable progress towards the goal is sometimes tied to the cost
of attaining~the goal on different schedules.
Cost is a vital input into two other types of analyses that also sometimes
have a role in standard setting. Cost is needed for a benefit-cost analysis
that addresses the economic efficiency of alternative regulations. Cost is
also an input into any analysis of the economic impact of each regulatory
alternative. An economic impact analysis deals with the consequences of
the regulation for small businesses, employment, prices, and market and
industry structure.
1.2 Purpose of Manual
The purpose of this Manual is two-fold: (1) to compile up-to-date capital
costs, operating, and maintenance expenses, and other costs for "add-on" air
pollution control systems and (2) to provide a comprehensive, concise, con-
1-2
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sistent, and easy-to-use procedure for estimating and (where appropriate)
escalating these costs. ("Add-on" systems are those installed downstream
of an air pollution source to control its emissions.)
The Manual estimating procedure rests on the notion of the "factored"
or "study" estimate, nominally accurate to within ± 30%. This type of
estimate is well suited to estimating control system costs intended for use
in regulatory development. Study estimates are sufficiently accurate, yet do
not require the detailed, site-specific data inputs needed to make definitive
or other more accurate types of estimates.
1.3 Organization of the Manual
This Manual is a revision of the 1987 edition of the EAB Control Cost
Manual, [I] which, in turn, was a revision of the edition completed in 1978.
This fourth edition of the Manual includes sizing and costing procedures
and data for electrostatic precipitators and updated cost data for ther-
mal and catalytic incinerators. This edition includes minor revisions to
Chapters 2 ("Cost Estimating Methodology"), 5 ("Fabric Filters"), and 6
("Electrostatic Precipitators") that serve to clarify the material therein.
In addition, Chapter 3 ("Thermal and Catalytic Incinerators") has been
rewritten to include device descriptions and sizing and costing procedures
that are clearer, more concise, and more representative. Chapter 4 ("Car-
bon Adsorbers") has also been revised to add a comprehensive example
problem and discussion on alternative methods for determining the carbon
working capacity.
As with the third edition, this edition has been issued in self-contained
chapters. Each chapter addresses a logically separate topic, which can
be either of a general nature (e.g., this introduction) or of a more spe-
cific, equipment-oriented nature (e.g., fabric filters). The chapters which
comprise this portion of the Manual are listed in Table 1.1, alongside the
sections in the 1987 Manual they replace.
As in the third edition, each type of equipment, background topic, etc.,
is given its own chapter number, for ease of identification and to reinforce
the intent that each chapter should stand alone. Also, each of the auxiliary
equipment items (e.g., ductwork) will be covered in a separate chapter.
1-3
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Table 1.1: Contents of the OAQPS Control Coat Manual (Fourth Edition)
No.
1
2
3
4
5
6
New Chapter
Title
"Introduction"
"Cost Estimating Methodology"
"Thermal and Catalytic Incinerators"
"Carbon Adsorbers"
"Fabric Filters"
"Electrostatic Precipitators"
No.
1
2
3
4
5
-
Old Section Replaced
Title
"Introduction"
"Manual Estimating
Methodology"
"Thermal and Catalytic
Incinerators"
"Carbon Adsorbers"
"Fabric Filters"
None
This has been done mainly to eliminate the confusion that arises when
classifying auxiliaries like mechanical collectors which can either support a
primary control device or be control devices in their own right.
Each of these stand-alone chapters contains a:
• Process description, where the types, uses, and operating modes of
the equipment item and (if applicable) its auxiliaries are discussed;
• Sizing procedure, which enables one to use the parameters of the pol-
lution source (e.g., gas volumetric flow rate) to size the equipment
item(s) in question;
• Capital and annual costs for the equipment and suggested factors to
use in estimating these costs from equipment design and operational
(e.g., operating hours) parameters. These costs are presented in both
graphical and equations! forms wherever possible.
1.4 Intended Users of the Manual
As explained in Section 1.2, the Manual provides comprehensive procedures
and data for sizing and costing control equipment. Some of these proce-
1-4
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dures are based on rigorous engineering principles, such as the material and
engineering balances in Chapter 3. To fully appreciate, and correctly apply,
these procedures the user should be able to understand them. Moreover, the
user should be able to exercise "engineering judgement" on those occasions
when the procedures may need to be modified or disregarded. Typically,
only engineers and others with strong technical backgrounds possess this
kind of knowledge. Hence, the Manual is oriented toward the technical not
the non-technical user.
1.5 "Uniqueness" of the Manual
The Manual presents a different perspective on estimating air pollution
control system costs than other cost-oriented reports, such as:
• The Cost Digest: Coat Summaries of Selected Environmental Control
Technologies [2]
• A Standard Procedure for Cost Analysis of Pollution Control Opera-
tions]^}
• Handbook: Control Technologies for Hazardous Air Pollutants [4]
Although these reports (as well as many of the NSPS Background In-
formation Documents) contain costs for add-on control systems, they do
not duplicate the Manual for one or more of the following reasons: (1) their
costs have been based either wholly or partly on data in the previous Man-
uals] (2) they apply to specific source categories only, whereas the Manual
data may be applied generally; (3) their estimating procedures and costs
are of less than study estimate quality; or (4) they are not intended for
estimating costs used in regulatory development.
Reason (3) applies to the Cost Digest, for example, as this report, de-
signed for use by non-technical personnel, contains procedures for mak-
ing order-of-magnitude estimates (± 30% accuracy or worse). A Stan-
dard Procedure, conversely, was primarily intended for estimating costs for
R&D cases (e.g., demonstration projects), where some site-specific data
are available. Further, although the latter report contains a thorough list
1-5
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of equipment installation factors, it contains few equipment cost*. The
report, Handbook: Control Technologies, used data and estimating proce-
dures from the 1978 Manual to provide sound generalized procedures for
estimating thermal and catalytic incinerator costs. The fourth edition of
the Manual updates and expands this information.
Finally, since its inception, the Manual has been extensively used in
Agency regulatory development efforts. Accordingly, the Manual's role in
the speciality of air pollution control system cost estimating is both unique
and secure.
1-6
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References
[1] EAB Control Coat Manual (Third Edition), EPA, Office of Air Quality
Planning and Standards, Economic Analysis Branch, February 1987
(EPA 450/5-87-001A [NTISPB 87-166583/AS]).
[2] DeWolf, Glenn, et al. (Radian, Inc.), The Cost Digest: Cost Sum-
maries of Selected Environmental Control Technologies, EPA, ORD,
Office of Environmental Engineering and Technology, October 1984
(EPA-600/884-010).
[3] Uhl, Vincent W., A Standard Procedure for Coat Analysis of Pollution
Control Operations, Volumes I and II, EPA, ORD, Industrial Environ-
mental Research Laboratory, June 1979 (EPA-600/8-79-018a).
[4] Handbook: Control Technologies for Hazardous Air Pollutants, EPA,
ORD, Air and Energy Engineering Research Laboratory, September
1986 (EPA-625/6-86-014).
1-7
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Chapter 2
COST ESTIMATING
METHODOLOGY
William M. Vatavuk
Standards Development Branch, OAQPS
U. S. Environmental Protection Agency
Research Triangle Park, NC 27711
November 1989
Contents
2.1 Types of Cost Estimates 2-3
2.2 Cost Categories Defined 2-5
2.2.1 Elements of Total Capital Investment 2-5
2.2.2 Elements of Total Annual Cost 2-8
2.3 Engineering Economy Concepts 2-9
2.3.1 Time Value of Money 2-9
2-1
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2.3.2 Cash Flow 2-11
2.3.3 Annualization and Discounting Methods 2-14
2.4 Estimating Procedure . 2-15
2.4.1 Facility Parameters and Regulatory Options 2-15
2.4.2 Control System Design 2-16
2.4.3 Sizing the Control System 2-18
2.4.4 Estimating Total Capital Investment 2-20
2.4.4.1 General Considerations 2-20
2.4.4.2 Retrofit Cost Considerations 2-22
2.4.5 Estimating Annual Costs 2-24
2.4.5.1 Raw Materials 2-24
2.4.5.2 Operating Labor 2-25
2.4.5.3 Maintenance 2-26
2.4.5.4 Utilities 2-26
2.4.5.5 Waste Treatment and Disposal 2-27
2.4.5.6 Replacement Parts 2-28
2.4.5.7 Overhead 2-28
2.4.5.8 Property Taxes, Insurance, and Administra-
tive Charges 2-29
2.4.5.9 Capital Recovery 2-29
References - 2-31
2-2
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This chapter presents a methodology that will enable the user, having
knowledge of the source being controlled, to produce study-level cost esti-
mates for a control system to control that source. The methodology, which
applies to each of the control systems included in this Manual, is gen-
eral enough to be used with other "add-on" systems as well. Further, the
methodology may also be applicable to estimating costs of fugitive emission
controls and of other non-stack abatement methods.
Before presenting this methodology in detail, we should first discuss
the various kinds of cost estimates and then define the cost categories and
engineering economy concepts employed in making the estimates.
2.1 Types of Cost Estimates
As noted above, the costs and estimating methodology in this Manual are
directed toward the "study" estimate, of ± 30% accuracy. According to
Perry's Chemical Engineer's Handbook, a study estimate is "...used to
estimate the economic feasibility of a project before expending significant
funds for piloting, marketing, land surveys, and acquisition ... [However] it
can be prepared at relatively low cost with minimum data."[l] Specifically,
to make a study estimate, the following must be known:
• Location of the source within the plant;
• Rough sketch of the process flow sheet (i.e., the relative locations of
the equipment in the system);
• Preliminary sizes of, and material specifications for, the system equip-
ment items;
• Approximate sizes and types of construction of any buildings required
to house the control system;
• Rough estimates of utility requirements (e.g., electricity);
• Preliminary flow sheet and specifications for ducting and piping;
• Approximate sizes of motors required.[1]
2-3
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In addition, an estimate of the labor hours required for engineering and
drafting is needed, as the accuracy of an estimate (study or otherwise)
is highly dependent on the amount of engineering work expended on the
project.
There are, however, four other types of estimates, three of which are
more accurate than the study estimate. These are:[l]
• Order-of-magnitude. This estimate provides "a rule-of-thumb proce-
dure applied only to repetitive types of plant installations for which
there exists good cost history". Its error bounds are greater than ±
30%. (However, according to Perry's, "...no limits of accuracy can
safely be applied to it.") The sole input required for making this
level of estimate is the control system's capacity (often measured by
the maximum volumetric flow rate of the gas passing through the
system). So-called "six-tenths factor" estimates (not to be confused
with factored estimates) are examples of this type.
• Scope or Budget authorization or Preliminary. This estimate, nomi-
nally of ± 20% accuracy, requires more detailed knowledge than the
study estimate regarding the site, flow sheet, equipment, buildings,
etc. In addition, rough specifications for the insulation and instru-
mentation are also needed.
• Project control or Definitive. These estimates, accurate to within ±
10%, require yet more information than the scope estimates, espe-
cially concerning the site, equipment, and electrical requirements.
• Firm or Contractor's or Detailed. This is the most accurate (± 5%) of
the estimate types, requiring complete drawings, specifications, and
site surveys. Further, "[tjime seldom permits the preparation of such
estimates prior to an approval to proceed with the project."[l]
For the purposes of regulatory development, study estimates have been
found to be acceptable, as they represent a compromise between the less
accurate order-of-magnitude and the more accurate estimate types. The
former are too imprecise to be of much value, while the latter are not
only very expensive to make, but require detailed site and process-specific
knowledge that most Manual users will not have available to them.
2-4
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2.2 Cost Categories Defined
The names given certain categories of costs and what they contain vary
considerably throughout the literature. Certain words like "capital cost"
can have vastly different meanings, which can often lead to confusion, even
among cost estimators. To avoid this confusion and, at the same time,
provide uniformity in the Manual, basic terms are defined in this chapter
and will be used throughout. The terminology used is adapted from that
of the American Association of Cost Engineers.[2] Although it has been
developed for general use, it is readily adaptable to air pollution control
system costing.
First, two general kinds of costs are estimated, total capital investment
(TCI) and total annual coat (TAG). These are discussed below.
2.2.1 Elements of Total Capital Investment
The total capital investment includes all costs required to purchase equip-
ment needed for the control system (termed purchased equipment costs), the
costs of labor and materials for installing that equipment (termed direct in-
stallation costs), costs for site preparation and buildings, and certain other
costs which are termed indirect installation costs. The TCI also includes
costs for land, working capital, and off-site facilities.
Direct installation costs include costs for foundations and supports,
erecting and handling the equipment, electrical work, piping, insulation,
and painting. Indirect installation costs include such costs as engineering
costs; construction and field expenses (i.e., costs for construction supervi-
sory personnel, office personnel, rental of temporary offices, etc.); contractor
fees (for construction and engineering firms involved in the project); start-
up and performance test costs (to get the control system running and to
verify that it meets performance guarantees); and contingencies. Contin-
gencies is a catch-all category that covers unforeseen costs that may arise,
including (but certainly not limited to) "... possible redesign and modifi-
cation of equipment, escalation increases in cost of equipment, increases in
field labor costs, and delays encountered in start-up."[2]
2-5
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These elements of total capital investment are displayed in Figure 2.1.
Note that the sum of the purchased equipment cost, direct and indirect in-
stallation costs, site preparation, and buildings costs comprises the battery
limits estimate. By definition, this is the total estimate "... for a specific
job without regard to required supporting facilities which are assumed to
already exist... "[2] at the plant. This would mainly apply to control sys-
tems installed in existing plants, though it could also apply to those systems
installed in new plants when no special facilities for supporting the control
system (i.e., off-site facilities) would be required.
Where required, these off-site facilities would encompass units to pro-
duce steam, electricity, and treated water; laboratory buildings, railroad
spurs, roads, and the like. It is unusual, however, for a pollution control
system to have one of these units (e.g., a power plant) dedicated to it.
The system needs are rarely that great. However, it may be necessary—
especially in the case of control systems installed in new or "grass roots"
plants—for extra capacity to be built into the site generating plant to ser-
vice the system. (A venturi scrubber, which often requires large amounts of
electricity, is a good example of this.) It is customary for the utility costs
to be charged to the project as operating costs at a rate which covers both
the investment and operating costs for the utility.
As Figure 2.1 shows, there are two other costs which may be included
in the total capital investment for a control system. These are working
capital and land. The first of these, working capital, is a fund set aside
to cover the initial costs of fuel, chemicals, and other materials, as well as
labor and maintenance. It usually does not apply to control systems, for
the quantities of utilities, materials, labor, etc., they require are usually
small. (An exception might be an oil-fired thermal incinerator, where a
small supply (e.g., 30-day) of distillate fuel would have to be available
during its initial period of operation.)
Land may also be required. But, since most add-on control systems
take up very little space (a quarter-acre or less) this cost would be rel-
atively small. (Certain control systems, such as those used for flue gas
desulfurization, require larger quantities of land for the process equipment,
chemicals storage, and waste disposal.)
Note also in Figure 2.1 that the working capital and land are non-
depreciable expenses. In other words, these costs are "recovered" when the
control system reaches the end of its useful life (generally in 10 to 20 years).
2-6
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Figure 2.1: Elements of Total Capital Investment
* Primary Control
Device
• Auxiliary Equipment
(including ductwork)
• Instrumentation"
• Sales Taxes'*
• Freight0
Purcl
Equip
C<
lased
iment
>st
• Foundations
and
Supports
• Handling
and Erection
• Electrical
• Piping
• Insulation
• Painting
Dii
Instal
Co
Site
ect Preparation"'*
lation
st* Buildings*'
i
Total Direct
Cost
• Engineering
• Construction and
Field Expenses
• Contractor Fees
• Start-up
• Performance Test
• Contingencies
'indi
Instal
Co
Total I
C<
rect
lation
st»
ndirect
>st
Land*
Working
Capital'
"Battery Limits"
Cost
Off-Site
Facilities*
Total Non-depreciable
Investment
Total Depreciable
Investment
Total Capital
Investment
* Typically factored from the sum of the primary control device and auxiliary equipment costs.
* Typically factored from the purchased equipment cost.
" Usually required only at "grass roots" installations.
* Unlike the other direct and indirect costs, costs for these items usually are not factored from the
purchased equipment cost. Rather, they are sized and costed separately.
* Normally not required with add-on control systems.
2-7
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Conversely, the other capital costs are depreciable, in that they cannot be
recovered and are included in the calculation of income tax credits (if any)
and depreciation allowances, whenever income taxes are considered in a
cost analysis. (In the Manual methodology, however, income taxes are not
considered. See Section 2.3.)
Notice that when 100% of the system costs are depreciated, no salvage
value is taken for the system equipment at the conclusion of its useful life.
This is a reasonable assumption for add-on control systems, as most of the
equipment, which is designed for a specific source, cannot be used elsewhere
without modifications. Even if it were reusable, the cost of disassembling
the system into its components (i.e., "decommissioning cost") could be as
high (or higher) than the salvage value.
2.2.2 Elements of Total Annual Cost
The Total Annual Cost (TAG) for control systems is comprised of three
elements:' direct costs (DC), indirect coats (1C), and recovery credits (RC),
which are related by the following equation:
TAG = DC + IC~RC (2.1)
Clearly, the basis of these costs is one year, as this period allows for seasonal
variations in production (and emissions generation) and is directly usable
in profitability analyses. (See Section 2.3.)
Direct costs are those which tend to be proportional or partially propor-
tional to the quantity of exhaust gas processed by the control system per
unit time. These include costs for raw materials, utilities (steam, electricity,
process and cooling water, etc.), waste treatment and disposal, maintenance
materials, replacement parts, and operating, supervisory, and maintenance
labor. Of these direct costs, costs for raw materials, utilities, and waste
treatment and disposal are variable, in that they tend to be a direct func-
tion of the exhaust flow rate. That is, when the flow rate is at its maximum
rate, these costs are highest. Conversely, when the flow rate is zero, so are
the costs.
Semivariable direct costs are only partly dependent upon the exhaust
flow rate. These include all kinds of labor, maintenance materials, and
replacement parts. Although these costs are a function of the gas flow
2-8
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rate, they are not linear functions. Even while the control system is not
operating, some of the semivariable costs continue to be incurred.
Indirect, or "fixed", annual costs are those whose values are totally
independent of the exhaust flow rate and, in fact, would be incurred even
if the control system were shut down. They include such categories as
overhead, administrative charges, property taxes, insurance, and capital
recovery.
Finally, the direct and indirect annual costs are offset by recovery cred-
its, taken for materials or energy recovered by the control system, which
may be sold, recycled to the process, or reused elsewhere at the site. These
credits, in turn, must be offset by the costs necessary for their processing,
storage, transportation, and any other steps required to make the recov-
ered materials or energy reusable or resalable. Great care and judgement
must be exercised in assigning values to recovery credits, since materials
recovered may be of small quantity or of doubtful purity, resulting in their
having less value than virgin material.
The various annual costs and their interrelationships are displayed in
Figure 2.2. A more thorough description of these costs and how they may
be estimated is given in Section 2.4.
2.3 Engineering Economy Concepts
As mentioned previously, the estimating methodology presented in Section
2.4 rests upon the notion of the "factored" or "study" estimate. However,
there are other concepts central to the cost analyses which must be un-
derstood. These are (J.) the time value of money, (2) cash flow, and (3)
annualization.
2.3.1 Time Value of Money
The time value of money is based on the truism that "... a dollar now
is worth more than the prospect of a dollar. ..at some later date." [3] A
measure of this value is the interest rate which "... may be thought of as
the return obtainable by the productive investment of capital."[3]
2-9
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Figure 2.2: Elements of Total Annual Cost
• Raw materials
• Utilities
- Electricity
- Fuel
- Steam
- Water
— Compressed air
• Waste treatment/
disposal
• Labor
- Operating
— Supervisory
- Maintenance
• Maintenance materi-
als
• Replacement parts
Variable
Semivariable —
Direct
Coats
• Overhead
• Property taxes
• Insurance
• Administrative
charges
• Capital recovery
Indirect
Costs
Materials
Energy
Recovery
Credits
Total
Annual
Cost
2-10
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2.3.2 Cash Flow
During the lifetime of a project, various kinds of cash expenditures are made
and various incomes are received. The amounts and timing of these expen-
ditures and incomes constitute the cash flows for the project. In control
system costing it is normal to consider expenditures (negative cash flows)
and unusual to consider income (positive cash flows), except for product or
energy recovery income. By the simplifying convention recommended by
Grant, Ireson, and Leavenworth[3], each annual expenditure (or payment)
is considered to be incurred at the end of the year, even though the payment
will probably be made sometime during the year in question. (The error
introduced by this assumption is minimal, however.) Figure 2.3, which
shows three hypothetical cash flow diagrams, illustrates these end-of-year
payments. In these diagrams, P represents the capital investment, while the
A's denote the end-of-year annual payments. Note that in all diagrams, the
cash flows are in constant (real) dollars, meaning that they do not reflect
the effects of inflation. Also note that in the top diagram (I), the annual
payments are different for each year. (These represent the control system
annual costs (exclusive of capital recovery) described in Section 2.2.) In
reality, these payments would be different, as labor and maintenance re-
quirements, labor and utility costs, etc., would vary from year to year. A
generally upward trend in annual costs would be seen, however.
In diagram II, these fluctuating annual payments have been converted
to equal payments. This can be done by calculating the sum of the present
values of each of the annual payments shown in diagram I and annualizing
the total net present value to equivalent equal annual payments via a capital
recovery factor. (See discussion in the following paragraphs and in Section
2.3.3.) Alternatively, it is adequate to choose a value of A equal to the sum
of the direct and indirect annual costs estimated for the first year of the
project. This assumption is in keeping with the overall accuracy of study
estimates and allows for easier calculations.
Finally, notice diagram III. Here, the annual costs (A1) are again equal,
while the capital investment (P) is missing. Put simply, P has been incor-
porated into A1, so that A1 reflects not only the various annual costs but
the investment as well. This was done by introducing another term, the
capital recovery factor (CRF), defined as follows: "when multiplied by a
present debt or investment , [the CRF] gives the uniform end-of-year pay-
ment necessary to repay the debt or investment in n years with interest
2-11
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Figure 2.3: Hypothetical Cash Flow Diagrams"
I.
Year:
0123456789 10
Ail A2
i
A3
1
A4
i
A6
'
Aa
1
Ar
-
A8
.
A9
,
AID
,
II.
Year:
01234567
8 9 10
P
'
A
i
A
A
.
A
i
A
t
A
A
•
A
i
A
i
A
i
III.
Year:
0123456789 1(
A1
i
A1
1
A1
A1
A1
A1
i
A1
A1
A1
A1
i
'All Values Are Constant Year (Real) Dollars
2-12
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rate i."[3] The product of the CRF and the investment (P) is the capital
recovery cost (CRC):
CRC = CRF x P (2.2)
where
c« -(££?! (2'3)
Therefore, A1 is the sum of A and the CRC, or:
A1 = A + CRF x P (2.4)
In this context, n is the control system economic life, which, as stated
above, typically varies from 10 to 20 years. The interest rate (t) used in
this Manual is a pretax marginal rate of return on private investment of
10% (annual). This value, which could also be thought of as a "real private
rate of return", is used in most of the OAQPS cost analyses and is in
keeping with current OAQPS guidelines[4] and the Office of Management
and Budget recommendation for use in regulatory analyses. [5]
It may be helpful to illustrate the difference between real and nominal
interest rates. The mathematical relationship between them is straightfor-
ward:^]
(1 -Mn) = (1 + OU + >•) (2.5)
where
in, i = the annual nominal and real interest rates, re-
spectively
r = the annual inflation rate
Clearly, the real rate does not consider inflation and is in keeping with
the expression of annual costs in constant (i.e., real) dollars.
The above procedure using the pre-tax marginal (or real) rate of return
on private investment is the appropriate method for assessing the costs
from the perspective of the entity having to install the pollution control
equipment. For example, costs developed with the above procedure can
appropriately be used for answering questions concerning the market re-
sponse to regulation like price increases, quantity adjustments, and reduced
profitability.
2-13
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In an idealized economy with perfectly competitive and complete mar-
kets, this private cost and the social cost would be equal. However, in a
more realistic economy in which allocation of resources is distributed by
taxes, credit restrictions, and other market imperfections, the cost to soci-
ety is different than the private costs for capital expenditures. The costs to
society are the relevant costs for use in answering questions about economic
efficiency. For example, benefit cost analysis and cost-effectiveness analy-
ses should focus on cost to society, not just the cost to the entity facing
additional pollution control costs.
EPA has adopted a new approach, a two-stage approach, to discounting
for social costs. This new approach begins with the same capital recovery
costs (CRC) described above using the same 10% pre-tax marginal rate of
return on private investment. The second step of the two-stage approach
involves "discounting" both direct and indirect annual costs and CRC back
to an initial date ("year 0") using a consumption rate of interest of 3%.
(See Section 2.3.3 for an explanation of the discounting concept.) This
results in a relatively higher cost of capital from society's perspective than
from the perspective of the entity facing additional control cost. A detailed
explanation of this procedure and when it should be employed is beyond
the scope of this document. A fuller explanation is given in draft EPA
guidelines [6]. However it is mentioned here because the CRC and direct
and indirect annual costs are inputs to the two-stage procedure and must
be sufficiently itemized to allow use in the two-stage procedure.
2.3.3 Annualization and Discounting Methods
The above method of smoothing out the investment into equal end-of-
year payments, is termed the equivalent uniform annual cash flow (EUAC)
method.[3] In addition to its inherent simplicity, this method is very use-
ful when comparing the costs of two or more alternative control systems
(i.e., those which are designed to control the same source to an equivalent
degree). In fact, the EUAC's—or simply the total annual costs—of two
competing systems may be compared even if both the systems have differ-
ent economic lives, say 10 and 20 years. We recommend that the EUAC
method be used for estimating control costs unless particular circumstances
preclude its use.
Comparisons of systems with different economic lives cannot be made,
2-14
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however, using the other two annualization (i.e., profitability analysis)
methods—present worth (PW) and internal rate of return (IRR). The
present worth (or discounted cash flow) method involves the discounting of
all cash flows occurring after year 0 (i.e., the system startup date) back to
year 0. These cash flows are discounted by multiplying each by a discount
factor, (Y^sij where m is the number of years from year 0 to the year in
which the cash flow is incurred. The sum of these discounted cash flows
is then added to the capital investment to yield the present worth of the
project. The alternative having the highest present worth would be selected
(in control system costing this is usually a negative number). But when
comparing the present worths of alternative systems, the system lifetimes
must be equal for the comparison to be valid.[3]
The third annualization method, internal rate of return (IRR), is similar
to the present worth method, in that it involves the discounting of a series
of unequal cash flows. However, where with the PW method the interest
rate, i, is set beforehand, in the IRR method the interest rate is solved for
(usually via trial-and-error) after arbitrarily setting the PW to zero. When
comparing alternative systems, the one with the highest "IRR" (interest
rate) is selected.[3] But here again, the alternative systems compared must
have equal economic lives.
2.4 Estimating Procedure
The estimating procedure used in the Manual consists of five steps: (1)
obtaining the facility parameters and regulatory options for a given facility;
(2) roughing out the control system design; (3) sizing the control system
components; (4) estimating the costs of these individual components; and
(5) estimating the costs (capital and annual) of the entire system.
2.4.1 Facility Parameters and Regulatory Options
Obtaining the facility parameters and regulatory options involves not only
assembling the parameters of the air pollution source (i.e., the quantity,
temperature, and composition of the emission stream(s)), but also compil-
ing data for the facility's operation. (Table 2.1 lists examples of these.)
2-15
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Note that two kinds of facility parameters are identified—intensive and
extensive. The former are simply those variables whose values are indepen-
dent of quantity or dimensions—i.e., the extent of the system. Conversely,
extensive parameters encompass all size-dependent variables, such as the
gas volumetric flow rate.
Like the facility parameters, the regulatory options are usually speci-
fied by others. These options are ways to achieve a predetermined emission
limit. They range from no control to maximum control technically achiev-
able. The option provided will depend, firstly, on whether the emission
source is a stack (point source), a process leak (process fugitives source) or
an unenclosed or partly enclosed area, such as a storage pile (area fugitives
source). Stacks are normally controlled by "add-on" devices. As discussed
above, this Manual will deal primarily with these add-on devices. (How-
ever, some of these devices can be used to control process fugitives in certain
cases, such as a fabric filter used in conjunction with a building evacuation
system.) Add-ons are normally used to meet a specified emission level, al-
though in the case of particulate emissions, they may also be required to
meet an opacity level.
2.4.2 Control System Design
Step 2—roughing out the control system design—first involves deciding
what kinds of systems will be priced (a decision that will depend on the
pollutants to be controlled, exhaust gas stream conditions, and other fac-
tors), and what auxiliary equipment will be needed. When specifying the
auxiliary equipment, several questions need to be answered:
• What type of hood (if any) will be needed to capture the emissions
at the source?
• Will a fan be needed to convey the exhaust through the system?
• Is a cyclone or another pre-cleaner needed to condition the exhaust
before it enters the control device?
• Will the captured pollutants be disposed of or recycled? How will
this be done?
2-16
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Table 2.1: Facility Parameters and Regulatory Options
Facility Parameters
Intensive
— Facility status (new or existing, location)
— Gas characteristics (temperature, pressure, moisture content)
— Pollutant concentration(s) and/or particle size distribution
Extensive
— Facility capacity
- Facility life
— Exhaust gas flow rate
— Pollutant emission rate(s)
Regulatory Options
• No control
• "Add-on" devices
— Emission limits
— Opacity limits
• Process modifications
— Raw material changes
— Fuel substitution
• Others
— Coal desulfurization
2-17
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• Can the on-site utility capacity (e.g., electricity) accommodate the
added requirements of the control system?
The kinds of auxiliary equipment selected will depend on the answers
to these and other site-specific questions. However, regardless of the source
being controlled, each system will likely contain, along with the control
device itself, the following auxiliaries:
• Hood, or other means for capturing the exhaust;
• Ductwork, to convey the exhaust from the source to, through, and
from the control system;
• Fan system (fan, motor, starter, inlet/outlet dampers, etc.), to move
the exhaust through the system;
• Stack, for dispersing the cleaned gas into the atmosphere.
2.4.3 Sizing the Control System
Once the system components have been selected, they must be sized. Sizing
is probably the most critical step, because the assumptions made in this
step will more heavily influence the capital investment than any other.
Before discussing how to size equipment, we need to define the term. For
the purposes of this Manual, "sizing" is the calculation (or estimation) of
certain critical design parameters for a control device against which the
equipment cost of that device is most accurately correlated. For instance,
the equipment cost of an electrostatic precipitator (ESP) is most often
correlated with its collecting area. This, in turn, is a function of the exhaust
volumetric flow rate, the overall collection efficiency, and the empirically-
determined migration velocity, the ESP critical parameter. (Table 2.2 lists
examples of these parameters. For a full description of the ESP sizing
procedure, see Chapter 6.)
Also listed in Table 2.2 are general parameters which must also be spec-
ified before the purchased cost of the system equipment can be estimated.
Note that, unlike the control device parameters, these may apply to any
kind of control system. These parameters include materials of construction
2-18
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Table 2.2: Examples of Typical Control Device Parameters [11]
General
• Material of construction: e.g., carbon steel
• Insulated? Yes
• Economic life: 20 yr
• Redundancy0: none
.Device-Specific
• Gas-to-cloth ratio ("critical parameter"): 3.0 to 1
• Pressure drop: 6.0 in w.g. (inches water gauge)
• Construction: standard (vs. custom)
• Duty: continuous (vs. intermittent)
• Filter type: shaker
• Bag material: polyester, 16-oz.
"Refers to whether there are any extra equipment items installed (e.g., fans) to
function in case the basic item becomes inoperative, so as to avoid shutting down
the entire system.
2-19
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(which may range from carbon steel to various stainless steels to fiberglass-
reinforced polyester), presence or absence of insulation, and the economic
or useful life of the system. As indicated in Section 2.3.2, this last parame-
ter is required for estimating the annual capital recovery costs. The lifetime
not only varies according to the type of the control system, but with the
severity of the environment in which it is installed. (Representative values
for the system life and the other control device parameters will be presented
in those chapters of the Manual covering them.)
2.4.4 Estimating Total Capital Investment
2.4.4.1 General Considerations
The fourth step is estimating the purchased equipment coat of the control
system equipment. These costs are available from this Manual lor the most
commonly used add-on control devices and auxiliary equipment. Each type
of equipment is covered in a separate chapter. (See Table of Contents.)
Most of these costs, in turn, have been based on data obtained from
control equipment vendors. There are over one hundred of these firms,
many of whom fabricate and erect a variety of control systems.[8] They have
current price lists of their equipment, usually indexed by model designation.
If the items for which costs are requested are fabricated, "off-the-shelf"
equipment, then the vendor can provide a written quotation listing their
costs, model designations, date of quotation, estimated shipment date, and
other information. (See Figure 2.4 for a sample quotation.) Moreover,
the quote is usually "F.O.B." (free-on-board) the vendor, meaning that no
taxes, freight, or other charges are included. However, if the items are
not off-the-shelf, they must be custom fabricated or, in the case of very
large systems, constructed on-site. In such cases, the vendor can still give
quotations—but will likely take much longer to do so and may even charge
for this service, to recoup the labor and overhead expenses of his estimating
department.
As discussed in Section 2.2 in this Manual, the total capital investment
is factored from the purchased equipment cost, which in turn, is the sum
of the base equipment cost (control device plus auxiliaries), freight, instru-
mentation, and sales tax. The values of these installation factors depend
2-20
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Figure 2.4: Typical Vendor Quotation
QUOTATION
r
(NOTE: Company name and address h»ve been deleted.)
1
HAIL DROP I 12
U.S. EPA
RESEARCH TRIANGLE PARK
DURHAM. NC 27711
ATTN: MR. BILL VATAXUK
QUOTATION NO.
DMI
85S23382
9-23-85
VERBAL - BUDGET
J
Tlwnk you tor your Inquiry. W* ara pMuad to submit our quotation a* follow*:
ITEM t\ PREHEATER
MODEL 191-19 SIZE «9 IMPERVITE SHELL t TUBE HEAT EXCHANGER WITH
55.8 SQ. FT. OF HEAT TRANSFER AREA AND CODE STAMPED
ITEM 12 CONDENSER
$ 7.147.00 EA.
7.430.00 EA.
MODEL 191-19 SIZE 112 IMPERVITE SHELL * TUBE HEAT EXCHANGER WIT)
74.5 SQ. FT. OF HEAT TRANSFER AREA AND CODE STAMPED
APPROVAL DUG'S 2-3 WEEKS AFTER RECEIPT OF ORDER.
THIS QUOTATION IS IN CONFIRMATION OF OUR PHONE CONVERSATION OF
9/18/85.
6 tO 8
or Muwmo Afmowti
PMCM m F.0.8. >.N*t300«y«.
Untew otlMrwlM «tattd thn* prtct* art «ub(«ct to acmplanc* within X
-------
on the type of the control system installed and are, therefore, listed in the
individual Manual chapters dedicated to them.
The costs of freight, instrumentation, and sales tax are calculated dif-
ferently from the direct and indirect installation costs. These items are
factored also, but from the base equipment cost (F.O.B. the vendor(s)).
But unlike the installation factors, these factors are essentially equal for all
control systems. Values for these are as follows:
Cost Range Typical
Freight
Sales tax
Instrumentation
0.01 - 0.10
0 - 0.08
0.05 - 0.30
0.05
0.03
0.10
The range in freight costs reflects the distance between the vendor and the
site. The lower end is typical of major U.S. metropolitan areas, while the
latter would reflect freight charges to remote locations such as Alaska and
Hawaii.[7] The sales tax factors simply reflect the range of local and state
tax rates currently in effect in the U.S.[9]
The range of instrumentation factors is also quite large. For systems
requiring only simple continuous or manual control, the lower factor would
apply. However, if the control is intermittent and/or requires safety backup
instrumentation, the higher end of the range would be applicable.[7] Finally,
some "package" control systems (e.g., incinerators covered in Chapter 3)
have built-in controls, whose cost is included in the base equipment cost.
In those cases, the instrumentation factor to use would, of course, be zero.
2.4.4.2 Retrofit Cost Considerations
The installation factors listed elsewhere in the Manual apply primarily to
systems installed in new facilities. These factors must be adjusted whenever
a control system is sized for, and installed in (i.e, "retrofitted") an existing
facility. However, because the size and number of auxiliaries are usually
the same in a retrofit situation, the purchased equipment cost of the con-
trol system would probably not be different from the new plant purchased
cost. An exception is the ductwork cost, for in many retrofit situations
exceptionally long duct runs are required to tie the control system into the
existing process.
2-22
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Each retrofit installation is unique; therefore, no general factors can be
developed. Nonetheless, some general information can be given concerning
the kinds of system modifications one might expect in a retrofit:
1. Auxiliaries. Again, the most important component to consider is the
ductwork cost. In addition, to requiring very long duct runs, some
retrofits require extra tees, elbows, dampers, and other fittings.
2. Handling and Erection. Because of a "tight fit", special care may need
to be taken when unloading, transporting, and placing the equip-
ment. This cost could increase significantly if special means (e.g.,
helicopters) are needed to get the equipment on roofs or to other
inaccessible places.
3. Piping, Insulation, and Painting. Like ductwork, large amounts of
piping may be needed to tie in the control device to sources of pro-
cess and cooling water, steam, etc. Of course, the more piping and
ductwork required, the more insulation and painting will be needed.
4. Site Preparation. Unlike the other categories, this cost may actu-
ally decrease, for most of this work would have been done when the
original facility was built.
5. Off-Site Facilities. Conceivably, retrofit costs for this category could
be the largest. For example, if the control system requires large
amounts of electricity (e.g., a venturi scrubber), the source's power
plant may not be able to service it. In such cases, the source would
have to purchase the additional power from a public utility, expand
its power plant, or build another one. In any case, the cost of elec-
tricity supplied to that control system would likely be higher than if
the system were installed in a new source where adequate provision
for its electrical needs would have been made.
6. Engineering. Designing a control system to fit into an existing plant
normally requires extra engineering, especially when the system is
exceptionally large, heavy, or utility-consumptive. For the same rea-
sons, extra supervision may be needed when the installation work is
being done.
7. Lost Production. This cost is incurred whenever a retrofit control
system cannot be tied into the process during normally scheduled
2-23
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maintenance periods. Then, part or all of the process may have to
be temporarily shut down. The revenue lost during this shutdown
period is a bonafide retrofit expense.
8. Contingency. Due to the uncertain nature of many retrofit estimates,
the contingency (i.e., uncertainty) factor in the estimate should be
increased.
From the above points, it is apparent that some or most of these in-
stallation costs would increase in a retrofit situation. However, there may
be other cases where the retrofitted installation cost would be less than
the cost of installing the system in a new plant. This could occur when
one control device, say an ESP, is being replaced by a more efficient unit—
a baghouse, for example. The ductwork, stack, and other auxiliaries for
the ESP might be adequate for the new system, as perhaps would be the
support facilities (power plant, etc.).
2.4.5 Estimating Annual Costs
Determining the total annual cost is the last step in the estimating proce-
dure. As mentioned in Section 2.2 the TAG is comprised of three compo-
nents—direct and indirect annual costs and recovery credits. Unlike the
installation costs, which are factored from the purchased equipment cost,
annual cost items are usually computed from known data on the system
size and operating mode, as well as from the facility and control device
parameters.
Following is a more detailed discussion of the items comprising the total
annual cost. (Values/factors for these costs are also given in the chapters
for the individual devices.)
2.4.5.1 Raw Materials
Raw materials are generally not required with control systems. Exceptions
would be chemicals used in absorbers or venturi scrubbers as absorbents or
to neutralize acidic exhaust gases (e.g., hydrochloric acid). Chemicals may
also be required to treat wastewater discharged by scrubbers or absorbers
2-24
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before releasing it to surface waters. But, these costs are only considered
when a wastewater treatment system is exclusively dedicated to the control
system. In most cases, a pro-rata waste treatment charge is applied. (See
also discussion below on Waste Treatment and Disposal.)
Quantities of chemicals required are calculated via material balances,
with an extra 10 to 20% added for miscellaneous losses. Costs for chemicals
are available from the Chemical Marketing Reporter and similar publica-
tions.
2.4.5.2 Operating Labor
The amount of labor required for a system depends on its size, com-
plexity, level of automation, and operating mode (i.e., batch or continu-
ous). The labor is usually figured on an hours-per-shift basis. As a rule,
though, data showing explicit correlations between the labor requirement
and capacity are hard to obtain. One correlation found in the literature is
logarithmic:[10]
- = (-Y
LI \VJ
(2.6)
n\ \KI/
where
LI 5 LS = labor requirements for systems 1 and 2
Vi, V2 = capacities of systems 1 and 2 (as measured by
the gas flow rate, for instance)
y = 0.2 to 0.25 (typically)
The exponent in Equation 2.6 can vary considerably, however. Con-
versely, in many cases, the amount of operator labor required for a system
will be approximately the same regardless of its size.
A certain amount must be added to operating labor to cover supervi-
sory requirements. Fifteen per cent of the operating labor requirement is
representative. [11]
To obtain the annual labor cost, multiply the operating and supervisory
labor requirements by the respective wage rates (in $/hr) and the system
operating factor (number of hours per year the system is in operation). The
2-25
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wage rates also vary widely, depending upon the source category, geograph-
ical location, etc. These data are tabulated and periodically updated by
the U.S. Department of Labor, Bureau of Labor Statistics, in its Monthly
Labor Review and in other publications. Finally, note that these are base la-
bor rates, which do not include payroll and plant overhead. (See Overhead
discussion below.)
2.4.5.3 Maintenance
Maintenance labor is calculated in the same way as operating labor aind is
influenced by the same variables. The maintenance labor rate, however, is
normally higher than the operating labor rate, mainly because more skilled
personnel are required. A 10% wage rate premium is typical.[11]
Further, there are expenses for maintenance materials—oil, other lubri-
cants, duct tape, etc., and a host of small tools. Costs for these items can
be figured individually, but since they are normally so small, they are usu-
ally factored from the maintenance labor. Reference [10] suggests a factor
of 100% of the maintenance labor to cover the maintenance materials cost.
2.4.5.4 Utilities
This cost category covers many different items, ranging from electricity to
compressed air. Of these, only electricity is common to all control devices,
where fuel oil and natural gas are generally used only by incinerators; water
and water treatment, by venturi scrubbers, quenchers, and spray chambers;
steam, by carbon adsorbers; and compressed air, by pulse- jet fabric filters.
Techniques and factors for estimating utility costs for specific devices
are presented in their respective sections. However, because nearly every
system requires a fan to convey the exhaust gases to and through it, a gen-
eral expression for computing the fan electricity cost (Ce) is given here: [7]
0.746
6356
2-26
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where
Q = gas flow rate (actual ft3/min)
AP = pressure drop through system (inches of water, gauge)
(Values for AP are given in the chapters covering the
equipment items.)
s = specific gravity of gas relative to air (1.000, for all practi-
cal purposes)
9 = operating factor (hr/yr)
77 = combined fan and motor efficiency (usually 0.60 to 0.70)
pe = electricity cost ($/kwhr).
A similar expression can be developed for calculating pump motor electric-
ity requirements.
2.4.5.5 Waste Treatment and Disposal
Though often overlooked, there can be a significant cost associated with
treating and/or disposing of waste material captured by a control system
that neither can be sold nor recycled to the process.
Liquid waste streams, such as the effluent from a venturi scrubber, are
usually processed before being released to surface waters. The type and
extent of this processing will, of course, depend on the characteristics of
the effluent. For example, the waste can first be sent to one (or more)
clarifiers, for coagulation and removal of suspended solids. The precipitate
from the clarifier is then conveyed to a rotary filter, where most of the liquid
is removed. The resulting filter cake is then disposed of, via landfilling, for
example.
The annual cost of this treatment is relatively high—$1.00 to $2.00/thou-
sand gallons treated or more. [12] The solid waste disposal costs (via land-
filling, for example) typically would add another $20 to $30/ton disposed
of.[13] This, however, would not include transportation to the disposal site.
More information on these technologies and their costs is found in Refer-
ences [12] and [13].
2-27
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2.4.5.6 Replacement Parts
This cost is computed separately from maintenance, because it is a large
expenditure, incurred one or more times during the useful life of a control
system. This category includes such items as carbon (for carbon adsorbers),
bags (for fabric filters) and catalyst (for catalytic incinerators), along with
the labor for their installation.
The annual cost of the replacement materials is a function of the initial
parts cost, the parts replacement labor cost, the life of the parts, and the
interest rate, as follows:
(2.8)
where
CRCp = capital recovery cost of replacement parts ($/yr)
Cp = initial cost of replacement parts, including taxes and
freight ($)
cost of parts- replacement labor ($)
capital recovery factor (defined in Section 2.3).
In the Manual methodology, replacement parts are treated the same as
any other investment, in that they are also considered an expenditure that
must be amortized over a certain period. Also, the useful life of the parts
(typically 2 to 5 years) is generally less than the useful life of the rest of
the control system.
Replacement-part labor will vary, depending upon the amount of the
material, its workability, accessibility of the control device, and other fac-
tors.
2.4.5.7 . Overhead
This cost is easy to calculate, but often difficult to comprehend. Much of
the confusion surrounding overhead is due to the many different ways it is
computed and to the several costs it includes, some of which may appear
to be duplicative.
2-28
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There are, generally, two categories of overhead, payroll and plant. Pay-
roll overhead includes expenses directly associated with operating, super-
visory, and maintenance labor, such as: workmen's compensation, Social
Security and pension fund contributions, vacations, group insurance, and
other fringe benefits. Some of these are fixed costs (i.e., they must be paid
regardless of how many hours per year an employee works). Payroll over-
head is traditionally computed as a percentage of the total annual labor
cost (operating, supervisory, and maintenance).
Conversely, plant (or "factory") overhead account for expenses not nec-
essarily tied to the operation and maintenance of the control system, in-
cluding: plant protection, control laboratories, employee amenities, plant
lighting, parking areas, and landscaping. Some estimators compute plant
overhead by taking a percentage of all labor plus maintenance materials
[10], while others factor it from the total labor costs alone.[2]
For study estimates, it is sufficiently accurate to combine payroll and
plant overhead into a single indirect cost. This is done in this Manual Also,
overhead is factored from the sum of all labor (operating, supervisory, and
maintenance) plus maintenance materials, the approach recommended in
reference [10]. The factors recommended therein range from 50 to 70% [10]
An average value of 60% is used in this Manual.
2.4.5.8 Property Taxes, Insurance, and Administrative Charges
These three indirect operating costs are factored from the system total
capital investment, and typically comprise 1, 1, and 2% of it, respectively.
Taxes and insurance are self-explanatory. Administrative charges covers
sales, research and development, accounting, and other home office ex-
penses. (It should not be confused with plant overhead, however.) For
simplicity, the three items are usually combined into a single, 4% factor.
This value, incidentally, is standard in all OAQPS cost analyses.
2.4.5.9 Capital Recovery
As discussed in Section 2.3, the annualization method used in the Manual is
the equivalent uniform annualized cost method. Recall that the cornerstone
of this method is the capital recovery factor which, when multiplied by the
2-29
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total capital investment, yields the capital recovery cost. (See Equation
2.2.)
However, whenever there are parts in the control system that must be
replaced before the end of its useful life, Equation 2.2 must be adjusted, to
avoid double-counting.
That is:
CRC. = CRF. (TCI - (Cp + Cp,)] (2.9)
where
CRC, = capital recovery cost for control system ($/yr)
TCI = total capital investment for entire system ($)
CRF, = capital recovery factor for control system.
The term (Cp + Cpi) accounts for the cost of those parts (including taxes
and freight) that would be replaced during the useful life of the control
system and the labor for replacing them. Clearly, CRF, and CRFP will not
be equal unless the control system and replacement part lives are equal.
2-30
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References
[1] Perry, Robert H., and Chilton, Cecil H., Perry's Chemical Engineers'
Handbook (Fifth Edition), McGraw-Hill, New York, 1973, pp. 25-12 to
25-16.
[2] Humphries, K. K. and Katell, S., Basic Cost Engineering, Marcel
Dekker, New York, 1981, pp. 17-33.
[3] Grant, E.L., Ireson, W.G., and Leavenworth, R.S., Principles of En-
gineering Economy, Sixth Edition, John Wiley & Sons, New York,
1976. .
[4] EAB (OAQPS) Guideline Memo: "Interest Rates for Regulatory Im-
pact Analyses (RIA)", May 27, 1982.
[5] Regulatory Program of the U. S. Government, Appendix V, Office of
Management and Budget, April 1,1988 - March 31,1989.
[6] Scheraga, Joel D., Draft of "Supplemental Guidelines on Discounting
in the Preparation of Regulatory Impact Analyses",Office of Policy,
Planning and Evaluation, U. S. EPA, March, 1989.
[7] Vatavuk, W. M. and Neveril, R. B., "Estimating Costs of Air-Pollution
Control Systems-Part I: Parameters for Sizing Systems," Chemical
Engineering, October 6, 1980, pp. 165-168.
[8] Pollution Equipment News 1989 Buyer's Guide, Rimbach Publishing,
Pittsburgh, 1989.
[9] Internal Revenue Service, Form 1040, 1985.
[10] Peters, M. S. and Timmerhaus, K. D., Plant Design and Economics for
Chemical Engineers (Third Edition), McGraw-Hill, New York, 1980.
2-31
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[11] Vatavuk, W. M. and Neveril, R. B., "Estimating Costs of Air-Pollution
Control Systems-Part II: Factors for Estimating Capital and Operat-
ing Costs," Chemical Engineering, November 3, 1980, pp. 157-162.
[12] Vatavuk, W. M. and Neveril, R. B., "Estimating Costs of Air-Pollution
Control Systems-Part XVII: Particle Emissions Control," Chemical
Engineering, April 2, 1984, pp. 97-99.
[13] The RCRA Risk-Coat Analysis Model, U.S. Environmental Protection
Agency, Office of Solid Waste, January 13, 1984.
2-32
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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
3-1
-------
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-H
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 Acid 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
3-2
-------
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 3.54
3.1 Introduction
Incineration, like carbon adsorption, is one of the best known methods of
industrial gas waste disposal. Unlike carbon adsorption, however, inciner-
ation 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
orgajiic_compound to be a VpjOjmlej£in8_s^ecificanyrdetermined not to^
j3eja_yQ£L4adeed, 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.
3-3
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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 mix-
ture. Combustion of such a mixture of organic compounds containing car-
bon, hydrogen, and oxygen is described by the overall exothermic reaction:
CXEV0Z + z + - 02 => xC02 + H20 (3.1)
L 4 21 2
The complete combustion products COj and EfoO are relatively innocu-
ous, making incineration an attractive waste disposal method. When chlo-
rijiajej^r^sulfj^cojtita|mn are present in the mTxtureTTKe
products of complete combustion include the acid components HCl_or SQ2,
respectivelyTth addition to H2O and CO^ In general, these streams^ would
require Te^novai-^fjfa^acid^^mj>o^ents by" a~scrubber~untt7^whlch~^could
greJatljTaffect The cost of the incineration system. (The sizing and cost-
l5g~oT these scrub berTTscovered in the "Wet Scrubbers" chapter of this
Manual.)
The heart' of an incinerator 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 com-
bustion, energy must be supplied to the incinerator 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 sufficient 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
incinerator or in a catalytic incinerator. In the catalytic incinerator a cat-
alyst 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 efficiency
3-4
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improved, by providing heat exchange between selected inlet streams and
the effluent stream. The effluent stream containing the products of combus-
tion, 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:
rac ion Energy actually recovered from flue gas . .
r> 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 in-
crease 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 efficiency using a "primary" heat exchanger.
Energy efficiency can be further improved by placing another ("sec-
ondary") exchanger downstream of the primary exchanger to recover addi-
tional 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 specific 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 deficient 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 incinerator. 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 stoichio-
metric mixture, presents an unacceptable fire hazard as a feed stream to
the incinerator. The lower, or fuel-lean, explosive limit (LEL) of a given
3-5
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organic compound defines the minimum concentration of that compound
in air that can produce more energy than is needed to raise its own temper-
ature 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 their LEL have a
heating value of approximately 50 Btu/scf.
Since the majority of industrial waste gases that contain VOCs are di-
lute mixtures of combustible gases in air, their heating value is low and
their oxygen content exceeds that required to combust both the waste or-
ganics (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 sat-
isfy fire insurance regulations. Generally, the streams are brought to below
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 below 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 incineration system is to control the amount of VOCs
released to the environment. Performance of a control device such as an
incinerator can be described by a control efficiency, defined according to
the following equation:
Control Eff.,% = I™* ™Sa **** V°C ~ °Utlet ma3S rate VOC1 x 100 (3.3)
I Inlet mass rate VOC J v '
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.
3-6
-------
Auz Fuel
Auz Air Qa
Emission Source
DUution Air
•ft*-
[»»•
Combustion
Chamber
¥asti
Prehi
QT'
) Gas
sater
Secondary
Energy Recovery
Stack
Figure 3.1: Thermal Incinerator - General Case
Performance of an incinerator can also be measured solely by the outlet
VOC concentration, usually in ppmv.
There are a number of different incinerator designs. These designs can
be broadly classified as thermal systems and catalytic systems. Thermal
systems may be direct flame incinerators with no energy recovery, flame
incinerators with a recuperative heat exchanger, or regenerative systems
which operate in a cyclic 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 supplemen-
tal 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 com-
pounds and is usually determined empirically. It is the temperature at
3-7
-------
which the combustion reaction rate (and consequently the energy produc-
tion 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 sufficiently high level.
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 level 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 species 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 incinera-
tor vendors can provide guidelines based on their extensive experience. In
practice, most streams are mixtures of compounds, thereby further compli-
cating the prediction of this temperature. Other studies [2,3,4], which are
based on actual field test data, show that commercial incinerators 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 conclude that mixing is a critical factor in de-
termining the destruction efficiency. Even though it cannot be measured,
3-8
-------
Table 3.1: Theoretical Reactor Temperatures Required for 99.99 Percent
Destruction by Thermal Incineration 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 parame-
ters, such as temperature. The most feasible and efficient way to improve
the mixing in an incinerator is to adjust it after start-up. The 98% control
level discussed in the previous paragraph presumes such an adjustment.
Ultimately, once the unit is built, it is the responsibility of the user to
operate and maintain the incinerator to insure compliance with applicable
regulations.
3.2.1.1 Direct Flame Incinerators
Many configurations of thermal incinerators 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 destruc-
tion efficiency. The simplest example of such a system is the direct flame
incinerator. With reference to Figure 3.1, the direct flame incinerator 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 incinerator.
3-9
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3.2.1.2 Recuperative Incinerators
Recuperative incinerators have improved energy efficiency as a result of
placing heat exchangers in the hot outlet gas streams. With reference
to Figure 3.1, the recuperative incinerator 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) Consider-
able 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" incinerator. 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 efficiency 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 efficiencies. 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 reli-
ability than plate-to-plate units.[5] In any case, because most incinerators
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 fuel in the combustion chamber and only moderate preheat
temperatures are sought in practice (<1200°F).
Secondary Energy Recovery (Additional Waste Energy Recov-
ery) 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 pre-
heating 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
level energy is needed to realize these savings.
3-10
-------
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 incineration 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 fuel 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 fuel 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 associated with these high-performance heat ex-
changers and combustion chambers may be offset by the increased auxiliary
fuel savings to make such a system economical. The costs of these regen-
erative units will be given separately in the cost correlations presented in
Section 3.5. Regenerative incinerators are not packaged units but are field-
erected only. Accordingly, the costs given in Section 3.5 for regenerative
units are for field-erected units.
3-11
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AnxAir
Mode A
Stack
Carainle Paeldaf
1f««41nv fl*fl
Caraiuio Pftcktnc
CooUnf Ou
I
1
*-
I
1
Combuvtlon'— | 1
Ctwmb« —J •
| 1
1
Auz Pu«l
Stack
Source f
Canonic Paeldnf
CooUn«GM
tl
Comburtie
Chamber
?
Aux Fuel
Mode B
Figure 3.2: Regenerable-type Thermal Incinerator
3-12
-------
Auz Fuel
Auz Air Qa
F
Emission Source
Dilution Air
•+>
1^
Preheat
Chamber
"*i
Catalyst
Chamber
«',
Taste Gas
Preheater
"T.
Stack
Figure 3.3: Catalytic Incinerator
3.2.2 Catalytic Incinerators
Catalytic incinerators employ a bed of active material (catalyst) that facil-
itates 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 incinerator 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 incinerator's effluent or other process heat or
both (Figure 3.3). The preheated gas stream is then passed over the cat-
alyst bed. The chemical reaction (combustion) between the oxygen in the
gas stream and the gaseous pollutants takes place at the catalyst surface.
Catalytic incineration can, in principle, 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 species 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 pol-
3-13
-------
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 cat-
alysts 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 -widely used
as solvents and degreasers and are often the subject of concern in VOC con-
trol. 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 oxida-
tion catalysts. Nevertheless, their concentration may be sufficiently 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 velocity) needed to
obtain a given level 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 cat-
alysts in a specific reactor design. This table shows that the temperature
required for this level of conversion of different VOCs on a given catalyst
and of the same VOC on different catalysts can vary significantly.
Particulate matter, including dissolved minerals in aerosols, can rap»idly
blind the pores of catalysts and deactivate them over time. Because essen-
tially all the active surface of the catalyst is contained in relatively 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 incinerator should -be constant for optimal
operation. Large fluctuations in the flow rate will cause the conversion
3-14
-------
Table 3.2: Catalyst Temperatures Required for Oxidizing 80% of Inlet
VOC to C02, °F for Two Catalysts
Compound Temperature, °F
Pi - Honeycomb
acrolein
n-butanol
n-propylamine
toluene
n-butyric acid
1,1,1-trichloroethane
dimethyl sulfide
382
413
460
476
517
661
—
294
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 manu-
facturing process generating the waste gas stream. It may be uneconomical
to change the process for the sake of making the operation of the catalytic
incinerator feasible. In such cases, thermal incinerators (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-coritaining stream with the catalyst
serves to distinguish catalytic incineration systems. Both fixed-bed and
fluid-bed systems are used.
3.2.2.1 Fixed-Bed Catalytic Incinerators
Fixed-bed catalytic incinerators may use a monolith catalyst or a packed-
bed catalyst. Each of these is discussed below.
Monolith Catalyst Incinerators The most widespread method of con-
tacting the VOC-containing stream with the catalyst is the catalyst mono-
3-15
-------
lith. In this scheme the catalyst is a porous solid block containing par-
allel, non-intersecting channels aligned in the direction of the gas flow.
Monoliths offer the advantages of minimal attrition due to thermal expan-
sion/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 particles are supported either 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 particles 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 addi-
tional advantage of fluid-beds is a high bed-side heat transfer as compared
to a normal gas heat transfer coefficient. 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 permis-
sible temperatures in the catalyst bed. In these reactors the gas phase
temperature rise from gas inlet to gas outlet is low, depending on the ex-
tent of heat transfer through imbedded heat transfer surfaces. The catalyst
temperatures depend on the rate of reaction occurring at the catalyst sur-
face 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 either fixed-bed or monolithic catalysts. This is
due to the constant abrasion of the fluidized catalyst pellets, which helps
remove these particulates from the exterior of the catalysts in a continuous
manner.
3-16
-------
A disadvantage of a fluid-bed is the gradual loss of catalyst by attri-
tion. Attrition-resistant catalysts have been developed to overcome this
disadvantage.[11]
3.2.3 Other Considerations—(Packaged versus Field-
Erected Units, Auxiliary Equipment)
3.2.3.1 Packaged vs. Field-Erected Units
With the exception of regenerative incinerators, the equipment cost cor-
relations included in this chapter are for packaged units only. They are
not valid for field-erected units. For regenerative incinerators, the correla-
tions are valid for field-erected units only. Packaged units are units that
have been shop fabricated and contain all elements necessary for operation,
except for connection to facilities at the site, e.g., utilities. The elements
include the combustion chamber, preheater, instrumentation, fan, and the
necessary structural steel, piping, and electrical equipment. This equip-
ment 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 correlations in this chapter are valid to 50,000 scfm for packaged units,
except for fluid-bed units which are valid to 25,000 scfm.
Conversely, 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 field-erected units are typically higher
than those for equivalent-sized packaged units because the factors that im-
prove efficiency of shop-fabrication, such as uniform working environment,
availability of tools and equipment, and more efficient work scheduling, are
generally not available in the field.
3.2.3.2 Acid Gas Scrubbers
The final outlet stream of any incineration system may contain certain pol-
lutants that must be removed. The combustion of sulfur-containing com-
3-17
-------
pounds results in SO2, while chlorinated compounds yield C12 and HC1 in
the product stream. These acid 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 participate removal as
well as acid gas scrubbing. In most cases adding a scrubber significantly
increases the cost of the incineration 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 special materials of construction. This added expense is not in-
cluded 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 included in the cost, and its de-
sign is usually done by the incineration unit vendor. The cost correlations
presented in this chapter include units both with and without energy re-
covery. 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 included in this
chapter.
3.2.3.4 Other Auxiliary Equipment
Additional auxiliary equipment such as hoods, ductwork, precoolers, cy-
clones, fans, motors, and stacks are addressed separately in other chapiters
of this Manual.
3-18
-------
3.2.4 Technology Comparison
Both the thermal and catalytic incineration 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 incinerator will fulfill these require-
ments at the same cost. This section presents a first step toward deciding
how best to deal with VOC emission abatement using incinerators consider-
ing some qualitative factors pertinent to the types of incinerators 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 incinerators is presented in Ta-
ble 3.3. From the earlier discussions, the following factors relating to the
presence of contaminants should be considered by potential users [12]:
• The fouling of the catalyst in a catalytic system is a possibility. Poi-
sons to the system include heavy metals, phosphorous, sulfur and
most halogens, although catalysts have been developed that are chlo-
rine resistant.
• The possibility of process upsets that could release any of the above
poisons or cause fluctuations in the heating value to the incinerator
would favor a thermal system.
•JExcept for the No.2 grade, fuel oil should not be considered as auxil-
iary fuel 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 fac-
tor relates to relative energy efficiency of the various types of incinerators:
• Thermal units generally require more auxiliary fuel than catalytic
units and operate at temperatures that are roughly 1000°F higher.
This difference in fuel requirement increases as the heating value of
the waste stream decreases.
3-19
-------
Table 3.3: Principal VOC Incinerator Technologies
Thermal Systems
• Direct Flame Incinerator
• Recuperative Incinerator (Direct Flame with Recuperative Heat Ex-
changer)
• Regenerative Incinerator Operating in a Cyclic Mode
• Fixed-Bed
— Monolith
- Packed-Bed
• Fluid-Bed
Catalytic Systems
3-20
-------
In general, a trade-off exists between the higher capital costs of catalytic
incinerators and the higher operating costs of thermal incinerators. 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 En-
ergy Balances
In the sizing and costing of the incinerator and the calculation of the auxil-
iary fuel requirements, it is necessary to make material and energy balances
around the entire incinerator 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 incineration process is a steady-state process, the accumulation
term is zero and the equation becomes
In — Out +'Generation = 0
For majj balances it is useful to restrict the balances to be made on the
mass of each atomic species 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
include a generation term. Thus, the simplified equations are
In — Out = 0 , for steady-state mass balances (3-5)
In — Out + Generation = 0 , for steady-state energy balances (3.6)
3-21
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For the incineration 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,
PQCP(T - Tref)
where
Cp = heat capacity
T = temperature.
The reference temperature, Tre/, is often taken to be zero or the tem-
perature of a convenient stream, e.g., the inlet gas stream, in whatever
units T is in, so the Tr
-------
of interest are
• flue gas flow rate, upon which all the equipment cost correlations are
based.
• 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 VOCa 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 fuel and auxiliary combustion air are needed, more
complete mass and energy balances must be made.
The design procedure illustrated below 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 specifications of the incinerator 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
• Oxygen content
• Chemical composition of the combustibles
3-23
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• Inerts content
• 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 of combustibles in air.
• Particulate content—The participate content is important if catalytic
incinerators are to be costed. An upstream filter may suffice if partic-
ulate content is too high. Fluid-bed catalytic incinerators can tolerate
higher participate contents than fixed-bed catalytic incinerators.
The following parameters must be specified for the incinerator:
• Desired control efficiency - This efficiency should be based on require-
ments dictated by relevant state and federal regulations.
• Combustion chamber outlet temperature - This temperature may
also be based on requirements of a regulation or on recommendations
developed during regulatory development.
• Desired percent energy recovery - The desired percent energy recov-
ery should be the result of a process optimization in which costs of
incinerators with several different levels of energy recovery are esti-
mated and the minimum cost design selected. The tradeoff is between
the capital cost of the energy recovery equipment and the operating
(fuel) costs.
Specifications 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 sufficient oxygen in the waste gas to support the
combustion of the waste organics (including VOCs) and the auxiliary fuel,
if auxiliary fuel 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 below:
Air Content, Vol. % = 100.0 - --°-? x 100 - — x 100 (3.7)
106 10*
= 99.8%
Oxygen Content,% = Air Content x 0.209 (3.8)
= 20.86%
3-24
-------
Table 3.4: Specification of Sample Problem
Variable
Value
Preheater Inlet Waste Gas Vol Flow Rate, Qw.,
Preheater Inlet Waste Gas Temp., T^., °F
Composition
Benzene Content, ppmv
Methyl Chloride Content, ppmv
Air Content
Particulate Content
Moisture Content
Desired Control Efficiency,%
Desired Percent Energy Recovery, HR,%
scfm
20,000
100
1000
1000
Balance
Negligible
Negligible
98
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 associated 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, LELmix, is given by Grelecki [13] as
n
V
-i
(3.9)
where
Xj = volume fraction of combustible component i
LELj = lower explosive limits of combustible component
j (PPmv)
n = number of combustible components in mixture
3-25
-------
For the example case,
n
,- = (1, 000 + 1, 000) xl(T8 (3.10)
t=i
~a
= 2,OOOxlO
From standard references [14] or from Appendix 3A,
= 14,000 ppmv for benzene
LEI/MC = 82,500 ppmv for methyl chloride
LEL = \ 1>00° i 1>00° I"' rim
[2,000 x 14,000 "*" 2,000 x 82,500j ( '
= 23,938 ppmv
w T „. total combustible cone, in mixture
% LELmi. = - — - - x 100 (3.12)
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, sufficient 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 of the waste
gas stream, (-A/iCw), Btu/scf The energy content of the gas stream,
expressed in terms of the heat of combustion, is calculated as follows:
3-26
-------
where
(-A/ie, ) = heat of combustion of the waste stream (Btu/scf )
(-A/iCl) = volumetric heat of combustion of component i at
25°C(Btu/scf)
Xi = volume fraction of component i in the waste gas
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 incinerator 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/iCB ) = 3,475 Btu/scf for benzene
* = 705 Btu/scf for methyl chloride
The compositions specified earlier as ppmv are converted to volume
fractions as follows:
SB. = 1,000 ppmv xlO~6 = 10~3 for benzene
SMC = 1>000 ppmv xlO~6 = 10~3 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
(-A/O = (3,475)(lO-3)+(705)(lO-3) (3.15)
= 4.18 Btu/scf
Assuming the waste gas is principally air, with a molecular weight of
28.97 and a corresponding density of 0.0739 Ib/scf, the heat of combustion
per pound of incoming waste gas is
(-A/O = 56.6 Btu/lb
3-27
-------
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 below 25%
LEL, which corresponds to about 13 Btu/scf. However, if LEL sensors and
monitors are installed, one can incinerate 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 excessively 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 (-AACllI), the design procedure for ther-
mal and catalytic incinerators is discussed separately, beginning with Step
5 for each type of incinerator.
3.4.2 Steps Specific to Thermal Units
Figure 3.1 shows a generic thermal incinerator 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 level of VOC
destruction. In general, state and local regulations specify the required level
of destruction that the customer must meet. In this example a destruction
efficiency of 98 percent is specified. Studies by Mascone [2,3,4] show that
this destruction efficiency 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 efficiency level is the minimum achievable by any new
properly designed and operated incinerator. Many incinerators can achieve
destruction efficiencies of 99% or higher.)
3-28
-------
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, T^, increases, the auxiliary fuel requirement decreases,
but at the expense of a larger heat exchanger. However, there are several
important Hmits on T^. First, Tw. must not be close to the ignition tem-
perature 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 acid-forming
atoms), the flue gas temperature after the heat exchanger, T/0, must not
drop below the acid dew. Both limitations limit the amount of heat ex-
change and thus the maximum value of TWo. 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 acid gases will result
in corrosion of many of the metals used in heat exchangers. As an example,
fuel sulfur contents of 1 to 2 percent can give acid 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 acid
dew elevation.
With the following assumptions, one can estimate TWo using equation
3.2, the definition of fractional energy recovery for a heat exchanger.
• The fractional energy recovery is specified.
• The amount of auxiliary fuel, Q0/, and auxiliary combustion air, Qa,
are small relative to the waste gas, Qu,, so that the mass flow rates
of gases, pwQw and p/Qf, on both sides of the preheater are approx-
imately the same, or
PwQw * PfQf
• The heat capacities of the gases on both sides of the preheater are
approximately the same, regardless of composition. This is true for
waste streams which are dilute mixtures of organics in air, the prop-
erties of the streams changing only slightly on combustion.
• The mean heat capacities above the reference temperature of the gases
on both sides of the preheater are approximately the same regardless
of temperature.
3-29
-------
With these assumptions, the equation for fractional energy recovery for a
heat exchanger becomes
T — T
Fractional Energy Recovery = -^ — — — (3.16)
For thi's example with a fractional energy recovery of 0.70, an incinerator
operating temperature, T/.., of 1,600°F, and a waste gas inlet temperature,
TWi, of 100°F, the waste gas temperature at the exit of the preheater be-
comes
Tw. = 1,150°F
The temperature of the exhaust gas, T/e, 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 capacities of the gases involved, results in the following equation:
Tft - Tf, = Tw, - Tw.
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
Tfa = 550°F
This value of T/0 should be well above the acid dew of the flue gas stream.
It should be remembered that Tw, should be well below the ignition tem-
perature of the VOC stream to prevent unwanted temperature excursions
in the preheater. This must be verified even if the stream is well below the
LEL because flammability limits can be expanded by raising the reactant
stream temperature. A sufficiently high preheat temperature, TWo, could
initiate reaction (with heat release) in the preheater. This would ordinarily
be detrimental to the materials of construction in the heat exchanger. The
one exception is the thermal incinerator of the regenerable type described
3-30
-------
in Section 3.2 The 95 percent energy recovery obtainable in regenerable
systems would result in this example in a TWt 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 sufficient 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 Q0/ using a mass and
energy balance around the combustion chamber and following the principles
discussed in Section 3.3, with reference to Figure 3.1.
• The reference temperature, Tre/, is taken as the inlet temperature of
the auxiliary fuel, Ta/.
• No auxiliary air, Q0, is required.
• Energy losses, EL, are assumed to be 10% of the total energy input to
the incinerator above ambient conditions.[16,17] Thus, if the reference
temperature is near ambient conditions,
HL = 0.1pfiQfiCpmfi(Tfi-Tref) (3.17)
• The heat capacities of the waste gases entering and leaving the com-
bustion chamber are approximately the same, regardless of compo-
sition. This is true for waste streams which are dilute mixtures of
organics in air, the properties of the streams changing only slightly
on combustion.
• The mean heat capacities above the reference temperature of the
waste gases entering and leaving the combustion chamber are ap-
proximately the same regardless of temperature. Thus the mean heat
capacity for the waste gas stream entering or leaving the combustion
chamber should be evaluated at the average of TWa and T/,.. For air
3-31
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this assumption introduces an error of, at most, 5% over the temper-
atures of interest.
With these assumptions, the mass and energy balance around the com-
bustion chamber reduces to the following equation:
ti - TWo - O.irre/) - (-
T
- lref)
Input data for this equation are summarized below:
• The waste stream is essentially air so that
p». = pwt = 0.0739 Ib/scf, air at 77°F, 1 atm.
Cpmair = 0.255 Btu/lb °F, the mean heat capacity of air between
77°F and 1,375°F (the average tem-
perature of the waste gas entering and
leaving the combustion chamber)
• Other input data to Equation 3.18 include
Qw. = Qm = 20, 000 scfm
(-A/iCa/) = 21,502 Btu/lb, for methane
Tref = Taf = 77°F, assume ambient conditions
paf = 0.0408 lb/ft3, methane at 77°F, 1 atm.
Tfi = 1,600°F, from Step 5t
TWo = 1,150°F, from Step 6t
(-A/iCwJ = 56.6 Btu/lb, from Step 4
Substituting the above values into Equation 3.18 results in
Qaf = 167 scfm
The values of the parameters in the energy balance are summarized in
Table 3.5.
3-32
-------
Table 3.5: Summary of Example Problem Variable Valuation
TPe/ = 77°F
x~\ f] rri
Stream Subscript, j Ib/scf scfm Btu/lb °F °F
IN - Sensible Heat
Auxiliary Air a na* na* na* na*
Auxiliary Fuel af 0.0408 167 ** 77
Waste Gas w0 0.0739 20,000 0.255 1,150
OUT - Sensible Heat
Waste Stream /.• 0.0739 20,167 0.255 1,600
(-Afce), waste gas = 56.6 Btu/lb
(-A/ie), 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 sensible 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 relatively 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 relatively 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 efficiency. In fact, with zero energy recovery in the sample problem,
the auxiliary fuel requirements would be 605 scfm, about four times the
energy requirements based on 70% energy recovery.
3-33
-------
Table 3.6: Terms in Energy Balance around Combustor - Example Prob-
lem
IN - OUT + GENERATION = 0
Stream _ Subscript, j Btu/min
IN - Sensible Heat, PjQjCpnt.(Ti - Tref) ~~
Auxiliary Air a 0
Waste Gas w0 404,403
OUT - Sensible Heat, PjQjCpnt.(Ti - !*«/)
Waste Stream ' ft 578,796
OUT - Losses
10% of total energy input 57,880
GENERATION -
Heat of Combustion, pjQj(—AhCi)
Waste Gas w0 83,655
Auxiliary Fuel af 146,506
3-34
-------
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 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 7t is sufficient for stabilization. If it is insufficient, then a
minimum amount of auxiliary fuel must be used, and the amount of energy
recovery specified earlier must be reduced to avoid exceeding the specified
temperature at which the incinerator will operate (Step 5t).
This check is made by calculating 5% of the total energy input to the
incinerator and comparing it with the auxiliary fuel energy input. The total
energy input is given as follows:
Total Energy Input = PfiQfiCpnji(Tfi - Tref) (3.19)
Auxiliary Fuel Energy Input = />„/
-------
(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 incinerator with the appropriate streams
labeled. The approach used in the calculations on the catalytic incinera-
tor is somewhat different than that used in the thermal incinerator. This
difference arises because of additional constraints which are placed on the
catalytic incinerator. These constraints are as follows:
• 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.
• The inlet temperature to the catalyst bed itself must be above the
catalytic ignition temperature required to give the desired destruction
efficiency in the incinerator. 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 desired reactor
bed outlet temperature.
• Auxiliary fuel, in combination with the preheat from the primary
heat exchanger, is used to preheat the waste gas to the reactor in-
let temperature. A minimum amount of auxiliary fuel (< 5% of the
total energy input) must be used to stabilize the flame in the pre-
heat combustion chamber. This has the effect of further limiting the
combustible content of the waste gas stream and the amount of heat
exchange permissible in the primary heat exchanger.
The steps outlined below represent one approach to recognizing these
constraints and incorporating them into the calculation procedures.
Step 5c - Establish the desired outlet temperature of the cata-
lyst bed, T/j The energy released by the oxidation of the VOCs in the
3-36
-------
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 specified that will ensure the
desired level of destruction of the VOC stream. As in thermal incinerators,
this value varies from compound to compound and also from catalyst to
catalyst. Final design of the incinerator should be done by firms with expe-
rience in incinerator 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 max-
imum bed temperature of 1,200°F should not be exceeded. In the example
problem the catalyst outlet temperature, TA, is selected to be 900°F.
Step 0c - 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 incinerator. The equation for fractional energy recovery,
Equation 3.16, is used, with the same assumptions as used for the thermal
incinerator. For the example problem with a fractional energy recovery of
0.70, a catalyst bed outlet temperature, T/rf, of 900°F, and a waste gas
inlet temperature, T^., of 100°F, the gas.temperature at the exit of the
preheater becomes
Tw, = 660°F
The same considerations regarding the closeness of the temperature of
the exhaust gas, T/e, to its dew point apply to the catalytic incinerator as
they did to the thermal incinerator.
Step 7c - Calculate the auxiliary fuel requirement, Qa/ The aux-
iliary fuel requirement, Qa/, is calculated by making mass and energy bal-
ances around the preheater combustion chamber and the catalyst chamber.
The auxiliary fuel requirement calculated in this manner must be checked to
insure that it falls within the constraints imposed by design considerations
of the catalytic incinerator. These constraints are as follows:
• The auxiliary fuel requirement must be positive. A negative fuel
requirement indicates that the heat of combustion of the waste gas,
3-37
-------
(—Afoc), is too high for the fractional energy recovery in the primary
heat exchanger that was selected.
• The auxiliary fuel 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 equa-
tion used in the thermal incinerator calculations. The input data for Equa-
tion 3.18 for the catalytic incinerator example problem are summarized
below:
• The waste stream is essentially air so that
Pw. = pwi = 0.0739 Ib/scf, air at 77°F, 1 atm
Cpma.r = 0.248 Btu/lb °F, the mean heat capacity of air be-
tween 77°F and 780°F(the average
of the preheater exit and catalyst
bed outlet temperatures)
• Other input data to Equation 3.18 include
Qw0 = Qu,i = 20,000 scfm
(-A/ic.,) = 21,502 Btu/lb, for methane
Taf = 77°F, assume ambient conditions
paf = 0.0408 lb/ft3, methane at 77°F, 1 atm
Tf. = 900°F, from Step 5c
TWa = 660°F, from Step 6c
(-A/ieJ = 56.6 Btu/lb, from Step 4
Substituting the above values into Equation 3.18 results in
Qaf = 40 scfm
If the outlet temperature of the catalyst bed, T/;, is 800°F, then Q0/
decreases to -6.7 scfm. In other words, no auxiliary fuel would, theoreti-
cally, be required at this bed temperature. However, as discussed above in
3-38
-------
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, (-A/ic,J, of about 79.9 Btu/lb would
cause the auxiliary fuel requirement, Q0/, 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 (-&hema) of
52.7 Btu/lb. Both of these heats of combustion are relatively 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 fuel (< 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 sufficient for stabilization. If it is insufficient, then a
minimum amount of auxiliary fuel must be used and the amount of energy
recovery specified earlier must be reduced to avoid exceeding the specified
temperature at which the incinerator will operate (Step 5c).
This check is made in the same manner as that in Step 8t of the thermal
incinerator calculation. The results of'this check indicate that the auxiliary
fuel requirement is more than sufficient to stabilize the burner flame.
Step 9c - Estimate the inlet temperature to the catalyst bed, Tri
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 selected for use.
The inlet temperature to the catalyst bed, Tri, should be such that,
when the temperature rise through the catalyst bed, AT, is added to it,
the resulting temperature is T/., 900°F. Thus,
AT = Tfi - Tri (3.21)
3-39
-------
The value of AT is determined by an energy balance around the pre-
heater portion of the combustor. The preheater is required to heat the
gases up to the catalyst bed inlet temperature using auxiliary fuel.1 This
energy balance is made with the assumptions made earlier in deriving Equa-
tion 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 /j subscript become terms with r^ sub-
scripts to denote a catalytic reactor inlet stream rather than a combustor
outlet (flue 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:
'. - T- ~
This equation may be rearranged to solve for Tr,. explicitly. This pro-
duces an equation that is somewhat complex and non-intuitive.
_ ref] + PV.QV.C^TV. + o.irre/)
l.lCp^>a/
-------
Step lOc - Calculate the total volumetric flow pate of gas through
the incinerator, Q/. The total volumetric flow rate of gas leaving the
incinerator is referred to as the flue gas flow rate, Q/{, and is the gas rate on
which the incinerator 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 num-
ber of moles of gas as a result of combustion is neglected, is the sum of the
inlet streams to the incinerator.
Qfi = Q*. + Q* + Q*f
= 20,000 + 0 + 40
= 20,040 scfm
Step He - Calculate the volume of catalyst in the catalyst bed
If the volumetric flow rate of gas through the catalyst bed, Q/n and the
nominal residence time (reciprocal space velocity) in the catalyst bed are
known, then the volume of catalyst can be estimated. There exists a com-
plex set of relationships between the catalyst volume and geometry, overall
pressure drop across the catalyst, conversion 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 being oxidized. It
is beyond the scope if this Manual to discuss these relationships, even in
an approximate way. For the purposes of cost estimation, the space ve-
locity, in reciprocal time units, necessary to achieve the required level of
destruction can be used to approximate the catalyst volume requirement.
The space velocity is defined as
where
Vcat = Overall bulk volume of the catalyst bed, including
interparticle voids (ft3)
By petro-chemical industry convention, the space velocity is computed at
the conditions of 60°F (not 77°F) and 1 atm. The volumetric flow rate,
3-41
-------
Q/t., must be corrected to these conditions. The proper space velocity to
achieve a desired level of conversion is based on experimental data for the
system involved. For precious metal monolithic catalysts, the space velocity
generally lies between 10,000 h"1 and 60,000 h"1. (Base metal catalysts
operate at lower space velocities, ranging from 5,000 to 15,000 h~1.)[10]
For the example, using a space velocity of 30,000 h"1 or 500 min"1, and
using QA at 60°F,
"' ~
= 19,400ft3/min
19,400ft3/min
500
= 39ft3
There are a number of catalyst bed parameters, such as catalyst configu-
ration and bed design, that are not significant 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 incineratorsr(l) thermal-recuperative, (2) thermal-
regenerative, (3) fixed-bed catalytic, and (4) fluid-bed catalytic.
3.5.1 Estimating Total Capital Investment
Total capital investment, TCI, includes the equipment cost, EC, for the
incinerator itself, the cost of auxiliary equipment (e.g., ductwork), all direct
and indirect installation costs, and costs for buildings, site preparation, off-
site facilities, land, and working capital. However, the last five costs usually
do not apply to incinerators. (See Chapter 2 of this Manual for a detailed
description of the elements comprising the TCI.)
3-42
-------
Table 3.7: Scope of Cost Correlations
Total (Flue) Gas
Incinerator Type Flowrate, scfm Figure Number
Thermal - Recuperative 500°-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 covers the 1,000 to 50,000 scfm range, the correlation
is valid for the 500 to 50,000 scfm range.
3.5.1.1 Equipment Costs, EC
As discussed in Section 3.2.3, the equipment costs, EC, given in this chap-
ter apply to packaged incinerators, except for regenerative incinerators.
For regenerative incinerators, the costs apply to field-erected units. The
EC typically includes all flange-to-flange equipment needed to oxidize the
waste gas, including the auxiliary burners, combustion chamber, catalyst,
primary heat exchanger (except for the "zero heat recovery" cases), weath-
ertight 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 incineration systems are presented in four groups delin-
eated according to their similarity of design. These groups are outlined in
Table 3.7. With the exception of regenerative thermal and fluid-bed cat-
alytic incinerators, 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 incinerator 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 inciner-
ation technology, design and manufacturing procedures vary from vendor
to vendor, so that costs may vary. As always, once the study estimate is
3-43
-------
completed, it is recommended that more than one vendor be solicited for a
more detailed cost estimate.
The additional expense of acid gas clean-up or particulate control is not
treated in this section. The equipment cost of a gas absorber to remove any
acid gases formed in the incinerator can be quite large, sometimes exceeding
the equipment cost of the incinerator itself even for simple packed tower
scrubbers [19]. For more complex absorbers that include 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 acid
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 ven-
dors [12,20,21]. The EC of these units are given as a function of total volu-
metric throughput, Qtot, in scfm. "Qt<,t" is the total volume of the gaseous
compounds exiting the combustion chamber; it is identical to the term,
"Q/j", used in Figures 3.1 and 3.3. This includes the combustion products,
nitrogen, unburned fuel 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 analysis, a log-log relationship between
throughput and EC was found for a given level of energy recovery (HR)
over the flow rate range from 500 to 50,000 scfm. These relationships are
as follows:
EC = 10294Q^355 HR = 0 % (3.24)
EC = 13149Q*f °9 HR = 35 % (3,25)
EC = 17056Q?J502 HR = 50 % (3.26)
EC = 21342g^500 HR = 70 % (3.27)
The regenerative (or excess enthalpy) systems provide up to 95 percent
heat recovery at the expense of higher capital costs. Their 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 approximately
3-44
-------
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linear function of total flow rate over a 10,000 to 100,000 scfm range by the
following equation:
EC = 2.204 x 10s + 11.57 Qtol (3.28)
Again, the higher capital costs of these units can be substantially offset
by the substantial savings in auxiliary fuel costs.
Catalytic Incinerators The EC for a catalytic incinerator 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 relationships with Qtot for each of the energy recovery rates.
Least squares regressions of the data yielded the following correlations for
total flow rates between 2,000 and 50,000 scfm:
EC = 1105Q™471 HR = 0% (3.29)
EC = 3623C&4189 HR = 35% (3.30)
EC = 1215<&Jm HR = 50% (3.31)
EC = 1443(&f57 HR = 70% (3.32)
Fluid-bed catalytic incinerators 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 species. For this en-
hanced flexibility of feed streams, a higher capital cost is incurred, as in-
dicated by the EC shown in Figure 3.7. The data shown were provided
by vendors [11,19] and exhibited a linear relationship 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 + 13.2QM HR = 0 % (3.33)
EC = 8.84 x 104 + 14.6Qtot HR = 35 % (3.34)
EC = 8.66 x 104 + 15.8<9tot HR = 50 % (3.35)
EC = 8.39 x 104 -I- 19.2Qtot 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
-------
EQUIPMENT COST, FOB, APRIL 1988 DOLLARS (THOUSANDS)
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PMENT COST, FOB, APRIL 1988 DOLLARS (THOUSANDS)
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A 50 % "
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FLUE GAS VOLUMETRIC FLOW RATE, SCFM (THOUSANDS)
Figure 3.7: Equipment Costs of Catalytic Incinerators, Fluid-Bed
3-49
-------
00
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O CATALYTIC - FLUID BED
D CATALYTIC - FIXED BED
A THERMAL - RECUP.
50% ENERGY RECOVERY
4 5 6 7 8 9 10
20
30 40 50 60 708090
FLUE GAS VOLUMETRIC FLOW RATE, SCFM (THOUSANDS)
Figure 3.8: Equipment Cost Comparison of Incinerator Types
3-50
-------
3.5.1.2 Installation Costs
As explained in Chapter 2, the purchased equipment cost, PEC, is calcu-
lated by taking the sum of the EC and the cost of auxiliary equipment
(e.j., 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 incinerators.
Table 3.9 shows the itemized installation costs that are obtained when
these installation factors are applied to the PECs for the example inciner-
ators. Depending on the site conditions, the installation costs for a given
incinerator could deviate significantly from costs generated by these aver-
age factors. Vatavuk and Neveril [25] provide some guidelines for adjusting
the average installation factors to account for other-than-average installa-
tion conditions. For units handling total gas flow rates lower than 20,000
scfm the installation costs are minimal, amounting normally to only util-
ity tie-ins (electrical and, if necessary, combustion or dilution air). The
installation costs for these smaller incinerators would be 20 to 25 % of the
>PEC. Smaller units may be installed on the roofs of manufacturing build-
ings rather than at ground level. 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 (TAG) is the sum of the direct and indirect annual
costs. The TAG 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 incinerators include labor (operating and supervi-
sory), maintenance (labor and materials), fuel, electricity, and (in catalytic
units) replacement catalyst. For thermal and catalytic units, the fuel usage
rate is calculated as shown in Sections 3.4.2 and 3.4.3, respectively, where
3-51
-------
Table 3.8: Capital Cost Factors for Thermal and Catalytic Incinerators"
Cost Item Factor
Direct Costs
Purchased equipment costs
Incinerator (EC) + auxiliary equipment* As estimated, A
Instrumentation* 0.10 A
Sales taxes 0.03 A
Freight 0.05 A
Purchased equipment cost, PEC B = 1.18 A
Direct installation costs
Foundations & supports 0.08 B
Handling & erection " 0.14 B
Electrical 0.04 B
Piping 0.02 B
Insulation for ductwork1* 0.01 B
Painting 0.01 B
Direct installation cost " 0.30 B
Site preparation As required, SP
Buildings As required, Bldg.
Total Direct Cost, DC 1.30 B 4- SP + Bldg.
Indirect Costs (installation)
Engineering 0.10 B
Construction and field expenses 0.05 B
Contractor fees 0.10 B
Start-up 0.02 B
Performance test 0.01 B
Contingencies 0.03 B
Total Indirect Cost, 1C Q31~B
Total Capital Investment = DC + 1C 1.61 B + SP + Bldg.
•Reference [25].
'Ductwork and any other equipment normally not included with unit furnished by incinerator vendor.
'Instrumentation and controls often furnished with the incinerator, and thus often included in the EC.
'if ductwork dimensions have been established, cost may be estimated based on SlO to 112/ft3 of
surface for field application. Fan housings and stacks may also be insulated.
3-52
-------
Table 3.9: Capital Costs for Thermal and Catalytic Incinerators
Example Problem
Cost Item
Direct Costs
Purchased equipment costs
Incinerator (EC)
Auxiliary equipment0
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
Site preparation0
Buildings0
Total Direct Cost
Indirect Costs (installation)
Engineering, 0.1 OB
Construction and field expenses, 0.05B
Contractor fees, 0.1 OB
Start-up, 0.02B
Performance test, 0.01B
Contingencies, 0.03B
Total Indirect Cost
Total Capital Investment (rounded)
Cost,
Thermal-
Recuperative
$254,200
—
$254,200
25,400
7,630
12,700
$300,000
24,000
42,000
12,000
6,000
3,000
3,000
$90,000
•^—
$390,000
30,000
15,000
30,000
6,000
3,000
9,000
$93,000
$483,000
$
Fluid-Bed
Catalytic
$468,200
—
$468,200
46,800
14,000
23,400
$552,466
44,200
77,300
22,100
11,000
5,520
5,520
$165,600
—
$718,000
55,200
27,600
55,200
11,000
5,520
16,600
$171,100
$889,000
'None of these items is required.
3-53
-------
Table 3.10: Annual Costs for Thermal and Catalytic Incinerators
Example Problem
Cost Item
Direct Annual Costs6, DC
Operating Labor
Operator
Supervisor
Operating materials
Maintenance
Labor
Material
Catalyst replacement
Utilities
Natural Gas
Electricity
Total DC
Indirect Annual Costs, 1C
Overhead
Administrative charges
Property taxes
Insurance
Capital recovery'
Total 1C
Total Annual Cost (rounded)
Suggested Factor Unit Cost'
0.5 h/shift $12.96/h
15% of operator —
—
0.5 h/shift $14.26/h
100% of —
maint. labor
100% of catalyst $650/ft3 for
replaced ea 2 yr metal oxide
— S3.30/kft3
— $0.069/kWh
60% of sum of —
operating, supv.,
& maint. labor k
maint. materials.
2% TCI —
1% TCI —
1% TCI —
CRP [TCI - —
1.08 (Cat. Cost)]
Thermal
6,480
972
7,130
7,130
0
264,500
35,000
9321,200
13,000
9,660
4,830
4,830
78,600
$110,900
$432,000
Fluid-Bed
Catalytic
6,480
972
7,130
7,130
14,600
63,400
42,300
8142,000
13,000
17,800
8,900
8,900
142,200
$190,800
$352,000
•1988 dollars.
'Assumes 8,000 h/yr.
cThe capital recovery cost factor, CRF, is a function of the catalyst or equipment life (typically, 2 and 10
years, respectively) and the opportunity cost of the capital (i.e., interest rate). For example, for a 10 year
equipment life and a 10% interest rate, CRF = 0.1628.
3-54
-------
Table 3.11: Typical Pressure Drop Across Selected Equipment
Equipment Type
Thermal Incinerators
Catalytic Fixed-bed Incinerators
Catalytic Fluid-bed Incinerators
Heat exchangers
» »
» n
Energy Recovery, %
0
0
0
35
50
70
A P, in.
4
6
6-10
4
8
15
H20
natural gas (methane) is assumed to be the fuel. (Other fuels could be used
for thermal units.)
The electricity costs are primarily associated with the fan needed to
move the gas through the incinerator. The power (in kilowatts) needed
to move a given volumetric flow rate of air (Qtot 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 efficiency, e, is adapted from Equation 2-7, as follows:
_ 1.17 x 10-4gtotAP
Power{m= -^ (3.37)
Fan efficiencies vary from 40 to 70 percent [15] while motor efficiencies are
usually 90 percent.
The total pressure drop across an incinerator system depends on the
number and types of equipment elements included in the system and on
design considerations. The estimation of actual pressure drop requirements
involves complex calculations based on the specific 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% efficiency for the fan and motor together. Using pressure drops of
4 and 8 inches of water, respectively, for the thermal and fluid-bed catalytic
incinerators2, 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
-------
heat recovery, the fan power requirements can be calculated as follows:
Thermal Incinerator
„ 1.17 x 10~4(20,000 scfm)(19 inches water)
PoWerfan = 1 _Ji 1
= 74.1 kW
Catalytic Incinerator
Power = 1'17 X 10"4(20'°00 scfm)(23 inches water)
°Werfan ~ 0.60
= 89.7 kW
The annual electricity costs would be the products of these usages, the
annual operating hours, and the electricity cost ($/kWh), or:
Electricity Cost (Thermal) = 74.1 kW x 8,000 hours/yr x $0.059/kWh
= $35,000 per yr
Electricity Cost (Catalytic) = 89.7 kW x 8,000 hours/yr x $0.059/kWh
= $42,300 per yr
The catalyst replacement costs and scheduling are highly variable and
depend on the nature of the catalyst, the amount of "poisons" and par-
ticulates in the gas stream (including the auxiliary fuel), the temperature
history of the catalyst, and the design of the unit. It is impossible to pre-
dict 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 field 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 com-
pound. 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
-------
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^ x 0.5762
= $14,600 per year
To calculate the fuel or electricity annual cost, multiply the fuel usage
rate (scfm) or the electricity 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., $/scfm for fuel and $/kWh for electricity).
For the example cases, the fuel costs can be calculated from the fuel
usage rates and the natural gas unit cost of $0.00330 /scf. For the thermal
incinerator example, the annual fuel cost is calculated as follows:
Annual Fuel n nnnnn $ „„,, scf ^min „ ^ hr
Tt, rr« i = 0-00330— x 167^- x 60-:— x 8,000—
, Inermal scf mm hr yr
= $264, 500 per yr
For the catalytic incinerator example, the annual fuel cost is found
similarly:
Annual Fuel 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 include 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, respectively. 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
-------
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 (Cent x 1.08 where the 1.08 is for freight 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 VOC 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 fuel
costs are significantly lower than those of the thermal unit. The disparity
is enough to offset the higher capital costs of the catalytic incinerator 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 level
of destruction met by the thermal incinerator 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 fuel requirements.
3.6 Acknowledgements
The authors gratefully acknowledge the following companies for contribut-
ing 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 Associates (EDA) (Itasca, IL)
Pillar Technologies, Inc. (Hartland, WI)
Huntington Energy Systems, Inc. (Union, NJ)
3-58
•
•
-------
• Regenerative Environmental Equipment Co. (REECO) (Morris Plains,
NJ)
• Englehard Corp. (Edison, NJ)
3-59
-------
-------
Appendix 3A
Properties of Selected
Compounds
3-60
-------
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
Cyclohexane
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
26.04
84.16
78.11
92.13
LEL°,
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
-------
Table 3.13: Molar Heat Capacities of Gases at Zero Pressure *
Cp = a + bT + cT* + dT3 ; T in °K
ffl CpdT
"• " (T, - TJ)
Cp in calories/g-mole °K or Btu/lb-mole °R
Compound
Methane
Ethane
Propane
Butane
Pentane
Hexane
Cyclopentane
Cyclohexane
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 xlO2
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 xlO5
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
-0.6861
0.3133
-0.4696
1.784
Temperature
Range, °K
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
-------
Table 3.14: Heats of Combustion of Selected Gaseous Organic Com-
pounds, —A/ic, at 25°C and constant pressure to form gaseous water and
carbon dioxide.*
Compound
Methane
Ethane
Propane
Butane
Pentane
Hexane
Octane
v Nonane
£ Decane
Ethylene
Propylene
Cyclopentane
Cyclohexane
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
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
10,476.7
9,698.4
9,784.7
'f*C
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 [15]
3-63
-------
References
[1] Prudent Practices for Disposal of Chemicals from Laboratories, Na-
tional Academy Press, Washington, D.C., 1983.
[2] Memo from Mascone, D.C., EPA, OAQPS, to Farmer, J. R., OAQPS,
EPA, June 11, 1980, Thermal Incinerator Performance for NSPS.
[3] Memo from Mascone, B.C., EPA, OAQPS, to Farmer, J. R., OAQPS,
EPA, July 22, 1980, Thermal Incinerator Performance for NSPS, Ad-
dendum.
[4] Memo from Mascone, D.C., EPA, OAQPS, to Farmer, J. R., OAQPS,
EPA, August 22,1980, Thermal Incinerators and Flares.
[5] Letter from Thomas Schmidt (ARI International, Palatine, IL) to
William M. Vatavuk (EPA, OAQPS, Research Triangle Park, NO),
August 16, 1989.
[6] Weldon, J. and S. M. Senkan, Combustion Sci. TechnoL, 1986, 47.
[7] Manning, P., Hazard Waste, 1984, 1(1).
[8] Pope, D., Walker, D. S., Moss, R. L., Atmos. Environ., 1976, 10.
[9] Musick, J. K., and F. W. Williams, Ind. Eng. Chem. Prod. Res. Dev.,
1974, 13(3).
[10] Letter from Robert M. Yarrington (Englehard Corporation, Edison,
NJ) to William M. Vatavuk (EPA, OAQPS, Research Triangle Park,
NC), August 14, 1989.
[11] Personal communication from Bill ShefFer (ARI, Inc., Palatine, II) to
Donald R. van der Vaart (RTI, Research Triangle Park, NC), March
30, 1988.
3-64
-------
[12] Personal communication from Ralph Stettenbenz (Combustion Engi-
neering, Air Preheater, Inc., Wellsville, NY) to Donald R. van der
Vaart (RTI, Research Triangle Park, NC), March 28, 1988.
[13] Grelecki, C., Fundamentals of Fire and Explosion Hazards Evaluation,
AIChE Today Series, New York, 1976.
[14] Weast, R. C. (ed.), CRC Handbook of Chemistry and Physics, 49th
ed., CRC Press, Cleveland, Ohio, 1968.
[15] Perry, R. H. and C. H. Chilton(eds.), Chemical Engineers Handbook,
5th ed., McGraw-Hill, New York, 1973.
[16] Personal Communication from Robert Yarrington (Englehard Corp.,
Edison, NJ) to William M. Vatavuk (EPA, OAQPS, Research Triangle
Park, NC), June 6, 1989.
[17] Personal Communication from Thomas Schmidt (ARI International,
Palatine, IL) to William M. Vatavuk (EPA, OAQPS, Research Triangle
Park, NC), June 7, 1989.
[18] Githens, R. E. and D. M. Sowards, Catalytic Oxidation of Hydrocarbon
Fumes, PB-299 132, National Technical Information Service, Spring-
field, VA.
[19] Personal Communication from Andrew Jones (Energy Development
Associates, Itasca, IL) to Donald R. van der Vaart (RTI, Research
Triangle Park, NC), March 4, 1988.
[20] Personal Communication from C. L. Bumford (Peabody Engineering,
Stamford, CT) to Donald R. van der Vaart (RTI, Research Triangle
Park, NC), March 28, 1988.
[21] Personal Communication from C. M. Martinson (TEC Systems, De-
Pere, WI) to Donald R. van der Vaart (RTI, Research Triangle Park,
NC), March 28, 1988.
[22] Personal Communication from Ronald J. Renko (Huntington Energy
Systems, Inc., Union, NJ) to Donald R. van der Vaart (RTI, Research
Triangle Park, NC), March 16, 1988.
[23] Personal Communication from James H. Mueller (Regenerative Envi-
ronmental Equipment Co., Inc., Morris Plains, NJ) to Donald R. van
der Vaart (RTI, Research Triangle Park, NC), January 13, 1988.
3-65
-------
[24] Personal Communication from Robert Hablewitz (Pillar Technologies,
Hartland, WI) to Donald R. van der Vaart (RTI, Research Triangle
Park, NC), March 20, 1988.
[25] Vatavuk, W. M. and R. Neveril, "Estimating Costs of Air Pollution
Control Systems, Part II: Factors for Estimating Capital and Operat-
ing Costs", Chemical Engineering, November 3, 1980, pp. 157-162.
[26] Kobe, K. A. and associates, "Thermochemistry for the Petrochemical
Industry", Petroleum Refiner, Jan. 1949 thru Nov. 1954.
3-66
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Chapter 4
CARBON ADSORBERS
William M. Vatavuk
Standards Development Branch, OAQPS
U. S. Environmental Protection Agency
Research Triangle Park, NC 27711
William L. Klotz
Chas. T. Main, Inc.
Charlotte, NC 28224
Robert L. S tailings
Research Triangle Institute
Research Triangle Park, NC 27709
November 1989
4-1
-------
-------
Contents
4.1 Process Description 4-3
4.1.1 Introduction 4-3
4.1.2 Types of Adsorbers 4-4
4.1.2.1 Fixed-bed Units 4-4
4.1.2.2 Cannister Units 4-7
4.1.3 Adsorption Theory 4-8
4.2 Design Procedure 4-14
4.2.1 Sizing Parameters 4-14
4.2.2 Determining Adsorption and Desorption Times .... 4-16
4.2.3 Estimating Carbon Requirement 4-18
4.2.3.1 Overview of Carbon Estimation Procedures . 4-18
4.2.3.2 Carbon Estimation Procedure Used in Manua/4-18
4.3 Estimating Total Capital Investment 4-20
4.3.1 Fixed-Bed Systems 4-20
4.3.1.1 Carbon Cost 4-21
4.3.1.2 Vessel Cost 4-21
4.3.1.3 Total Purchased Cost 4-24
4.3.1.4 Total Capital Investment 4-25
4.3.2 Cannister Systems 4-25
4.4 Estimating Total Annual Cost 4-27
4-2
-------
4.4.1 Direct Annual Costs 4-28
4.4.1.1 Steam 4-28
4.4.1.2 Cooling Water 4-28
4.4.1.3 Electricity 4-29
4.4.1.4 Carbon Replacement 4-32
4.4.1.5 Solid Waste disposal 4-32
4.4.1.6 Operating and Supervisory Labor 4-33
4.4.1.7 Maintenance Labor and Materials 4-33
4.4.2 Indirect Annual Costs 4-33
4.4.3 Recovery Credits 4-34
4.4.4 Total Annual Cost 4-35
4.4.5 Example Problem 4-35
References 4-43
4.1 Process Description
4.1.1 Introduction
In air pollution control, adsorption is employed to remove volatile organic
compounds (VOC's) from low to medium concentration gas streams, when
a stringent outlet concentration must be met and/or recovery of the VOC
is desired. Adsorption itself is a phenomenon where gas molecules passing
through a bed of solid particles are selectively held there by attractive forces
which are weaker and less specific than those of chemical bonds. During
adsorption, a gas molecule migrates from the gas stream to the surface
of the solid where it is held by physical attraction releasing energy—the
"heat of adsorption", which approximately equals the heat of condensation.
4-3
-------
Adsorptive capacity of the solid for the gas tends to increase with the
gas phase concentration, molecular weight, difFusivity, polarity, and boiling
point.
Some gases form actual chemical bonds with the adsorbent surface
groups. This phenomenon is termed "chemisorption".
Most gases ("adsorbates") can be removed ("desorbed") from the ad-
sorbent by heating to a sufficiently high temperature, usually via steam or
(increasingly) hot combustion gases, or by reducing the pressure to a suf-
ficiently low value (vacuum desorption). The physically adsorbed species
in the smallest pores of the solid and the chemisorbed species may re-
quire rather high temperatures to be removed, and for all practical pur-
poses cannot be desorbed during regeneration. For example,, approximately
^Jo^5_^ercenjt_oj[._orga^iics adsorbed on virgnT activated carbon is either
Chemisorbed or very strongly physically adsorbed and, for all intents, can-
not be desorHed during regeneration.[1]
Adsorbents in large scale use include activated carbon, silica gel, ac-
tivated alumina, synthetic zeolites, fuller's earth, and. other clays. This
chapter is oriented toward the use of activated carbon, a commonly used
adsorbent for VOCs.
4.1.2 Types of Adsorbers
Five types of adsorption equipment are used in collecting gases: (1) fixed
regenerable beds; (2) disposable/rechargeable cannisters; (3) traveling bed
adsorbers; (4) fluid bed adsorbers; and (5) chromatographic baghouses.[2]
Of these, the most commonly used in air pollution control are the fixed-bed
and cannister types. This chapter addresses only fixed-bed and cannister
units.
4.1.2.1 Fixed-bed Units
Fixed-bed units can be sized for controlling continuous, VOC-containing
streams over a wide range of flow rates, ranging from several hundred to
several hundred thousand cubic feet per minute (cfm). The VOC concen-
tration of streams that can be treated by fixed-bed adsorbers can be as
4-4
-------
low as several parts per billion by volume (ppbv) in the case of some toxic
chemicals or as high as 25% of the VOCs' lower explosive limit (LEL). (For
most VOCs, the LEL ranges from 2500 to 10,000 ppmv.[3|)
Fixed-bed adsorbers may be operated in either intermittent or contin-
uous modes. In intermittent operation, the adsorber removes VOC for a
specified time (the "adsorption time"), which corresponds to the time dur-
ing which the controlled source is emitting VOC. After the adsorber and
the source are shut down (e.g., overnight), the unit begins the deaorption
cycle during which the captured VOC is removed from the carbon. This
cycle, in turn, consists of three steps: (1) regeneration of the carbon by
heating, generally by blowing steam through the bed in the direction op-
posite to the gas flow;1 (2) drying of the bed, with compressed air or a fan;
and (3) cooling the bed to its operating temperature via a fan. (In most
designs, the same fan can be used both for bed drying and cooling.) At the
end of the desorption cycle (which usually lasts 1 to l| hours), the unit sits
idle until the source starts up again.
In continuous operation a regenerated carbon bed is always available for
adsorption, so that the controlled source can operate continuously without
shut down. For example, two carbon beds can be provided: while one is
adsorbing, the second is desorbing/idled. As each bed must be large enough
to handle the entire gas flow while adsorbing, twice as much carbon must
be provided than an intermittent system handling the same flow. If the
desorption cycle is significantly shorter than the adsorption cycle, it may
be more economical to have three, four, or even more beds operating in
the system. This can reduce the amount of extra carbon capacity needed
or provide some additional benefits, relative to maintaining a low VOC
content in the effluent. (See Section 4.2 for a more thorough discussion of
this.)
A typical two-bed, continuously operated adsorber system is shown in
Figure 4.1. One of the two beds is adsorbing at all times, while the other
is desorbing/idled. As shown here, the VOC-laden gas enters vessel #1
through valve A, passes through the carbon bed (shown by the shading)
and exits through valve B, from whence it passes to the stack. Meanwhile,
vessel #2 is in the deaorption cycle. Steam enters through valve C, flows
1 Although steam is the most commonly used regenerant, there are situations where it
should not be used. An example would be a degreasing operation that emits halogenated
VOCs. Steaming might cause the VOCs to decompose
4-5
-------
(Drying/Cooling I
Air) A
Waste Gas
(Fro*
Source)
Syste* Fan
(Drying/
Cooling Air)
ii
Vessel II
7{{77//////////////A
Stew—CXj-1
SteaM
Sttw-
VOC Vapor
Out In
(Cooling!
Hater I
Total
Condenser
VOC
Condensate •*•
(To Storage.
Processing)
1
Decanter
Hater
(To Treatment/
Sewer)
• (To Stack)
Figure 4.1: Typical Two-Bed, Continuously Operated Fixed-Bed Carbon
Adsorber System
4-6
-------
through the bed and exits through D. The steam-VOC vapor mixture
passes to a condenser, where cooling water condenses the entire mixture.
If part of the VOC is immiscible in water, the condensate next passes to a
decanter, where the VOC and water layers are separated. The VOC layer
is conveyed to storage. If impure, it may receive additional purification by
distillation. Depending on its quality (i.e., quantity of dissolved organics),
the water layer is usually discharged to a wastewater treatment facility.
Once steaming is completed, valves C and D are closed and valve E is
opened, to allow air to enter to dry and cool the bed. After this is done, the
bed is placed on standby until vessel #1 reaches the end of its adsorption
cycle. At this time, the VOC-laden gas is valved to vessel #2, while vessel
#1 begins its desorption cycle, and the above process is repeated.
In Figure 4.1, the system fan is shown installed ahead of the vessels,
though it could also be placed after them. Further, this figure does not
show the pumps needed to bring cooling water to the condenser. Nor does
it depict the solvent pump which conveys the VOC condensate to stor-
age. Also missing are preconditioning equipment used to cool, dehumidify,
or remove particulate from the inlet gases. Such equipment may or may
not be needed, depending on the condition of the inlet gas. In any case,
preconditioning equipment will not be covered in this chapter.
4.1.2.2 Cannister Units
Cannister-type adsorbers differ from fixed-bed units, in that they are nor-
mally limited to controlling low-volume, (typically 100 ft3/min, maximum)
intermittent gas streams, such as those emitted by storage tank vents, where
process economics dictate that either toll regeneration or throw-away can-
nisters are appropriate. The carbon cannisters are not intended for desorp-
tion on-site. However, the carbon may be regenerated at a central facility.
Once the carbon reaches a certain VOC content, the unit is shut down, re-
placed with another, and disposed of or regenerated by the central facility.
Each cannister unit consists of a vessel, activated carbon, inlet connection
and distributer leading to the carbon bed, and an outlet connection for
the purified gas stream.[4] In one design (Calgon's Ventsorb®), 150 Ibs of
carbon are installed on an 8-inch gravel bed, in a 55-gallon drum. The type
of carbon used depends on the nature of the VOC to be treated.
4-7
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In theory, a cannister unit would remain in service no longer than a
regenerable unit would stay in its adsorption cycle. Doing so would help
to insure the allowable outlet concentration from being exceeded. In real-
ity, however, poor operating practice may result in the cannister remaining
connected until the carbon is near or at saturation. This is because: (1)
the carbon (and often the vessel) will probably be disposed of, so there is
the temptation to operate it until the carbon is saturated; and (2) unlike
fixed-bed units, whose outlet VOC concentrations are usually monitored
continuously (via flame ionization detectors, typically), cannisters are usu-
ally not monitored. Thus, the user can only guess at the outlet loading,
and could tend to leave a unit in place longer.
4.1.3 Adsorption Theory
At equilibrium, the quantity of gas that is adsorbed on activated carbon is
a function of the adsorption temperature and pressure, the chemical species
being adsorbed, and the carbon characteristics, such as carbon particle size
and pore structure. For a given adsorbent-VOC combination at a given
temperature, an adsorption isotherm can be constructed which relates the
mass of adsorbate per unit weight of adsorbent ("equilibrium adsorptivity")
to the partial pressure of the VOC in the gas stream. The adsorptivity in-
creases with increasing VOC partial pressure and decreases with increasing
temperature.
A family of adsorption isotherms having the shape typical of adsorption
on activated carbon is plotted in Figure 4.2. This and other isotherms whose
shapes are convex upward throughout, are designated "Type I" isotherms.
The Freundlich isotherm, which can be fit to a portion of a Type I curve,
is commonly used in industrial design.[2]
we = kPm (4.1)
where
• we = equilibrium adsorptivity (Ib adsorbate/lb adsorbent)
P = partial pressure of VOC in gas stream (psia)
k,m = empirical parameters
The treatment of adsorption from gas mixtures is complex and beyond
4-8
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Jj
o
fc.
o
I/I
•o
o
tf»
Adsorfaate Partial Pressure (psia)
Figure 4.2: Type I Adsorption Isotherms For Hypothetical Adsorbate
4-9
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the scope of this chapter. Except where the VOC in these mixtures have
nearly identical adsorption isotherms, one VOC in a mixture will tend to
displace another on the carbon surface. Generally, VOCs with lower vapor
pressures will displace those with higher vapor pressures, resulting in the
former displacing the latter previously adsorbed. Thus, during the course
of the adsorption cycle the carbon's capacity for a higher vapor pressure
constituent decreases. This phenomenon should be considered when sizing
the adsorber. To be conservative, one would normally base the adsorption
cycle requirements on the least adsorbable component in a mixture and the
desorption cycle on the most adsorbable component.[1]
The equilibrium adsorptivity is the maximum amount of adsorbate the
carbon can hold at a given temperature and VOC partial pressure. In
actual control systems, however, the entire carbon bed is never allowed to
reach equilibrium. Instead, once the outlet concentration reaches a preset
limit (the "breakthrough concentration"), the adsorber is shut down for
desorption or (in the case of cannister units) replacement and disposal. At
the point where the vessel is shut down, the average bed VOC concentration
may only be 50% or less of the equilibrium concentration. That is, the
carbon bed may be at equilibrium ("saturated") at the gas inlet, but contain
only a small quantity of VOC near the outlet.
As Equation 4.1 indicates, the Freundlich isotherm is a power function '
that plots as a straight line on log-log paper. Conveniently, for the concen-
trations/partial pressures normally encountered in carbon adsorber opera-
tion, most VOC-activated carbon adsorption conforms to Equation 4.1. At
very low concentrations, typical of breakthrough concentrations, a linear
approximation (on arithmetic coordinates) to the Freundlich isotherm is
adequate. However, the Freundlich isotherm does not accurately represent
the isotherm at high gas concentrations and thus should be used with care
as such concentrations are approached.
Adsorptivity data for selected VOCs were obtained from Calgon Cor-
poration, a vendor of activated carbon.[5] The vendor presents adsorptivity
data in two forms: a set of graphs displaying equilibrium isotherms [5] and
as a modification of the Dubinin-Radushkevich (D-R) equation, a semi-
empirical equation that predicts the adsorptivity of a compound based on
its adsorption potential and polarizability.[6] In this Manual, the modified
D-R equation is referred to as the Calgon fifth-order polynomial. The data
displayed in the Calgon graphs [5] has been fit to the Freundlich equa-
4-10
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Table 4.1: Parameters for Selected Adsorption Isotherms*0
Isotherm
Adsorption Parameters
Adsorbate Temp. (°F) k m
(1) Benzene
(2) Chlorobenzene
(3) Cyclohexane
(4) Dichloroethane
(5) Phenol
(6) Trichloroethane
(7) Vinyl Chloride
(8) m-Xylene
(9) Acrylonitrile
(10) Acetone
(11) Toluene
77
77
100
77
104
77
100
77
77
100
100
77
0.597
1.05
0.508
0.976
0.855
1.06
0.200
0.708
0.527
0.935
0.412
0.551
0.176
0.188
0.210
0.281
0.153
OM61
0.477
0.113
0.0703
0.424
0.389
0.110
Range of
isotherm*
(psia)
0.0001-0.05
0.0001-0.01
0.0001-0.05
0.0001-0.04
0.0001-0.03
0.0001-0.04
0.0001-0.05
0.0001-0.001
0.001-0.05
0.0001-0.015
0.0001-0.05
0.0001-0.05
* Reference [5].
0 Each isotherm is of the form: we = kPm. (See text for definition of
terms.) Data are for adsorption on Calgon type "BPL" carbon.
* Equations should not be extrapolated outside these ranges.
tion. The resulting Freundlich parameters are shown in Table 4.1 for a
limited number of chemicals. The adsorbates listed include aromatics (e.g.,
benzene, toluene), chlorinated aliphatics (dichloroethane), and one ketone
(acetone). However, the list is far from all-inclusive.
Notice that a range of partial pressures is listed with each set of pa-
rameters, k and m. (Note: In one case (m-xylene) the isotherm was so
curvilinear that it had to be split into two parts, each with a different set
of parameters.) This is the range to which the parameters apply. Ex-
trapolation beyond.this range—especially at the high end—can introduce
inaccuracy to the calculated adsorptivity.
• But high-end extrapolation may not be necessary, as the following will
show. In most air pollution control applications, the system pressure is
4-11
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approximately one atmosphere (14.696 psia). The upper end of the partial
pressure ranges in Table 4.1 goes from 0.04 to 0.05 psia. According to
Dalton's Law, at a total system pressure of one atmosphere this corresponds
to an adsorbate concentration in the waste gas of 2,720 to 3,400 ppmv. Now,
as discussed in Section 4.1.2, the adsorbate concentration is usually kept at
25% of the lower explosive limit (LEL).2 For many VOCs, the LEL ranges
from 1 to 1.5 volume %, so that 25% of the LEL would be 0.25 to 0.375%
or 2,500 to 3,750 ppmv, which approximates the high end of the partial
pressure ranges in Table 4.1.
Finally, each set of parameters applies to a fixed adsorption tempera-
ture, ranging from 77° to 104°F. These temperatures reflect typical oper-
ating conditions, although adsorption can take place as low as 32°F and
even higher than 104°F. As the adsorption temperature increases to much
higher levels, however, the equilibrium adsorptivity decreases to such an
extent that VOC recovery by carbon adsorption may become economically
impractical.
The Calgon fifth-order polynomial is somewhat more accurate than the
Freundlich parameters from Table 4.1. The polynomial contains a temper-
ature parameter, and it allows one to estimate adsorption isotherms for
compounds not shown in Table 4.1 if pure component data are available.
The pure component data required are the saturation pressure, liquid molar
volume, and the refractive index. It is, however, somewhat more complex
to use than the Freundlich equation. The Calgon fifth-order polynomial is
as follows:
The mass loading, we, is calculated from
we = -^— x (Molecular Wt of Adsorbate) (4.2)
Kn
where
we = mass loading, i.e., equilibrium adsorptivity (Ib adsor-
bate per Ib carbon)
G = carbon loading at equilibrium (ft3 liquid adsorbate per
100 Ib carbon)
Vm = liquid molar volume of adsorbate (ft3 per Ib-mole).
'Although, Factory Mutual Insurance will reportedly permit operation at up to 50%
of the LEL, if proper VOC monitoring is used.
4-12
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The carbon loading, G, is calculated from the Calgon fifth-order polynomial
where
A3Y3
ABY*
(4.3)
= 1.71
Ax = -1.46 x ID"2
A2 = -1.65 x 10~3
A3 = -4.11 x 1(T4
A4 = +3.14 x 10-6
A5 = -6.75 x 10-7
and Y is calculated from several equations which follow.
The first step in calculating Y is to calculate x- This can be done by
calculating the adsorption potential, e:
c = 0.556 RT \n(Pt/Pi)
(4.4)
where
R = 0.730 (atm-ft3 per lb-mole-°R)
T = absolute temperature (°R)
P, = vapor pressure of adsorbate at the temperature T
(psia)
Pi = partial pressure of adsorbate (psia).
The x is calculated from:
X = e/(2.303/ZVm)
By substituting for e in the above equation, x can alternatively be calcu-
lated from:
X = 0.00890 (T/Vm) loglo(P./PO.
The next step in calculating Y is to calculate the relative polariziability, F.
T = 0i/00
4-13
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where
©i = polarizability of component i per unit volume, where
component i is the adsorbate
00 = polarizability of component o per unit volume, where
component o is the reference component, n-heptane.
For the adsorbate or the reference compound, using the appropriate refrac-
tive index of adsorbate, n, the polarizability is calculated from:
"1-1
Once x and F are known, Y can be calculated from:
(4.5)
Calgon also has a proprietary seventh-order form in which two addi-
tional coefficients are added to the Calgon fifth-order polynomial, but the
degree of fit reportedly is improved only modestly. [6] Additional sources of
isotherm data include the activated carbon vendors, handbooks (such as
Perry's Chemical Engineer's Handbook), and the literature.
4.2 Design Procedure
4.2.1 Sizing Parameters
Data received from adsorber vendors indicate that the size and purchase
cost of a fixed-bed or cannister carbon adsorber system primarily depend
on four parameters:
1. The volumetric flow of the VOC laden gas passing through the carbon
bed(s);
2. The inlet and outlet VOC mass loadings of the gas stream;
3. The adsorption time (i.e., the time a carbon bed remains on-line to
adsorb VOC before being taken off-line for desorption of the bed);
4-14
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4. The working capacity of the activated carbon.
In addition, the cost could also be affected by other stream conditions,
such as the presence/absence of excessive amounts of particulate, moisture,
or other substances which would require the use of extensive pretreatment
and/or corrosive-resistant construction materials.
The purchased cost depends to a large extent on the volumetric flow
(usually measured in actual ft3/min). The flow, in turn, determines the
size of the vessels housing the carbon, the capacities of the fan and motor
needed to convey the waste gas through the system, and the diameter of
the internal ducting.
Also important are the VOC inlet and outlet gas stream loadings, the
adsorption time, and the working capacity of the carbon. These variables
determine the amount and cost of carbon charged to the system initially
and, in turn, the cost of replacing that carbon after it is exhausted (typi-
cally, five years after startup). Moreover, the amount of the carbon charge
affects the size and cost of the auxiliary equipment (condenser, decanter,
bed drying/cooling fan), because the sizes of these items are tied to. the
amount of VOC removed by the bed. The amount of carbon also has a
bearing on the size and cost of the vessels.
A carbon adsorber vendor[7] supplied data that illustrate the depen-
dency of the equipment cost on the amount of the carbon charge. Costs
were obtained for fixed-bed adsorbers sized to handle three gas flow rates
ranging from 4,000 to 100,000 scfrn and to treat inlet VOC (toluene) con-
centrations of 500 and 5,000 ppm. Each adsorber was assumed to have an
eight hour adsorption time. As one might expect, the equipment costs for
units handling higher gas flow rates were higher than those handling lower
gas flow rates. Likewise, at each of the gas flow rates, the units sized to treat
the 5,000 ppm VOC streams had higher equipment costs than those sized
to treat the 500 ppm concentration. These cost differences ranged from 23
to 29% and averaged 27%. These higher costs were partly needed to pay for
the additional carbon required to treat the higher concentration streams.
But some of these higher costs were also needed for enlarging the adsorber
vessels to accommodate the additional carbon and for the added struc-
tural steel to support the larger vessels. Also, larger condensers, decanters,
cooling water pumps, etc., were necessary to treat the more concentrated
streams. (See Section 4.3.)
4-15
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The VOC inlet loading is set by the source parameters, while the outlet
loading is set by the VOC emission limit. (For example, in many states,
the average VOC outlet concentration from adsorbers may not exceed 25
ppm.)
4.2.2 Determining Adsorption and Desorption Times
The relative times for adsorption and desorption and the adsorber bed con-
figuration (i.e., whether single or multiple and series or parallel adsorption
beds are used) establish the adsorption/desorption cycle profile. The cycle
profile is important in determining carbon and vessel requirements and in
establishing desorption auxiliary equipment and utility requirements. An
example will illustrate. In the simplest case, an adsorber would be control-
ling a process which emits a relatively small amount of VOC intermittently
—say, during one 8-hour shift per day. During the remaining 16 hours the
system would either be desorbing or on standby. Properly sized, such a sys-
tem would only require a single bed, which would contain enough carbon
to treat eight hours worth of gas flow at the specified inlet concentration,
temperature, and pressure. Multiple beds, operating in parallel, would
be needed to treat large gas flows (>100,000 actual ft3/min, generally)[7],
as there are practical limits to the sizes to which adsorber vessels can be
built. But, regardless of whether a single bed or multiple beds were used,
the system would only be on-line for part of the day.
However, if the process were operating continuously (24 hours), an extra
carbon bed would have to be installed to provide adsorptive capacity during
the time the first bed is being regenerated. The amount of this extra capac-
ity must depend on the number of carbon beds that would be adsorbing at
any one time, the length of the adsorption period relative to the desorption
period, and whether the beds were operating in parallel or in series. If one
bed were adsorbing, a second would be needed to come on-line when the
first was shut down for desorption. In this case, 100% extra capacity would
be needed. Similarly, if five beds in parallel were operating in a staggered
adsorption cycle, only one extra bed would be needed and the extra capac-
ity would be 20% (i.e., 1/5)—provided, of course, that the adsorption time
were at least five times as long as the desorption time. The relationship
between adsorption time, desorption time, and the required extra capacity
4-16
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can be generalized.
Me = MClxf (4.6)
where
Mc, Mc/ = amounts of carbon required for continuous or in-
termittent control of a given source, respectively
(Ibs)
/ = extra capacity factor (dimensionless)
This equation shows the relationship between Me and Mc/. Section 4.2.3
shows how to calculate these quantities.
The factor, /, is related to the number of beds adsorbing (N^) .and
desorbing (Np) in a continuous system as follows:
(Note: N^ is also the number of beds in an intermittent system that would
be adsorbing at any given time. The total number of beds in the system
would be N4 + N/j.)
It can be shown that the number of desorbing beds required in a contin-
uous system (Np) is related to the desorption time (#0), adsorption time
(9 A), and the number of adsorbing beds, as follows:
A
(4.8)
(Note: BD is the total time needed for bed regeneration, drying, and cool-
ing.)
For instance, for an eight hour adsorption time, in a continuously operated
system of seven beds (six adsorbing, one desorbing) BD would have to be
l| hours or less (8 hours/6 beds). Otherwise, additional beds would have
to be added to provide sufficient extra capacity during desorption.
4-17
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4.2.3 Estimating Carbon Requirement
4.2.3.1 Overview of Carbon Estimation Procedures
Obtaining the carbon requirement (Mc or Me/) is not as straightforward as
determining the other adsorber design parameters. When estimating the
carbon charge, the sophistication of the approach used depends on the data
and calculational tools available.
One approach for obtaining, the carbon requirement is a rigorous one
which considers the unsteady-state energy and mass transfer phenomena
occurring in the adsorbent bed. Such a procedure necessarily involves a
number of assumptions in formulating and solving the problem. Such a
procedure is beyond the scope of this Manual at the present time, although
ongoing work in the Agency is addressing this approach.
In preparing this chapter of the Manual, we have adopted a rule-of-
thumb procedure for estimating the carbon requirement. This procedure,
while approximate in nature, appears to have the acceptance of vendors
and field personnel. It is sometimes employed by adsorber vendors to make
rough estimates of carbon requirement and is relatively simple and easy to
use. It normally yields results incorporating a safety margin, the size of
which depends on the bed depth (short beds would have less of a safety
margin than deep beds), the effectiveness of regeneration, the particular
adsorbate and the presence or absence of impurities in the stream being
treated.
4.2.3.2 Carbon Estimation Procedure Used in Manual
The rule-of-thumb carbon estimation procedure is based on the "working
capacity" (wci Ib VOC/lb carbon). This is the difference per unit mass
of carbon between the amount of VOC on the carbon at the end of the
adsorption cycle and the amount remaining on the carbon at the end of the
desorption cycle. It should not be confused with the "equilibrium capacity"
(we) defined above in section 4.1,3. Recall that the equilibrium capacity
measures the capacity of virgin activated carbon when the VOC has been
in contact with it (at a constant temperature and partial pressure) long
enough to reach equilibrium. In adsorber design, it would not be feasible
4-18
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to allow the bed to reach equilibrium. If it were, the outlet concentration
would rapidly increase beyond the allowable outlet (or "breakthrough")
concentration until the outlet concentration reached the inlet concentration.
During this period the adsorber would be violating the emission limit.
The working capacity is some fraction of the equilibrium capacity. Like
the equilibrium adsorptivity, the working capacity depends upon the tem-
perature, the VOC partial pressure, and the VOC composition. The work-
ing capacity also depends on the flow rate and the carbon bed parameters.
The working capacity, along with the adsorption time and VOC inlet
loading, is used to compute the carbon requirement for a cannister adsorber
or for an intermittently operated fixed-bed adsorber as follows:
(4.9)
where mvoc = VOC inlet loading (Ib/h)
Combining this with Equations 4.6 and 4.7 yields the general equation
for estimating the system total carbon charge for a continuously operated
system:
Me = I*2
-------
As Equation 4.10 shows, the carbon requirement is directly proportional
to the adsorption time. This would tend to indicate that a system could be
designed with a shorter adsorption time to minimize the carbon requirement
(and equipment cost). There is a trade-off here not readily apparent from
Equation 4.10, however. Certainly, a shorter adsorption time would require
less carbon. But, it would also mean that a carbon bed would have to be
desorbed more frequently. This would mean that the regeneration steam
would have to be supplied to the bed(s) more frequently to remove (in
the long run) the same amount of VOC. Further, each time the bed is
regenerated the steam supplied must heat the vessel and carbon, as well
as drive off the adsorbed VOC. And the bed must be dried and cooled
after each desorption, regardless of the amount of VOC removed. Thus,
if the bed is regenerated too frequently, the bed drying/cooling fan must
operate more often, increasing its power consumption. Also, more frequent
regeneration tends to shorten the carbon life. As a rule-of-thumb, the
optimum regeneration frequency for fixed-bed adsorbers treating streams
with moderate to high VOC inlet loadings is once every 8 to 12 hours.[1]
4.3 Estimating Total Capital Investment
Entirely different procedures should be used to estimate the purchased costs
of fixed-bed and cannister-type adsorbers. Therefore, they will be discussed
separately.
4.3.1 Fixed-Bed Systems
As indicated in the previous section, the purchased cost is a function of
the volumetric flow rate, VOC inlet and outlet loadings, the adsorption
time, and the working capacity of the activated carbon. As Figure 4.1
shows, the adsorber system is made up of several different items. Of these,
the adsorber vessels and the carbon comprise from one-half to nearly 90%
of the total equipment cost. (See Section 4.3.1.3.) There is also auxiliary
equipment, such as fans, pumps, condensers, decanters, and internal piping.
But because these usually comprise a small part of the total purchased
cost, they may be "factored" from "the costs of the carbon and vessels
without introducing significant error. The costs of these major items will
4-20
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be considered separately.
4.3.1.1 Carbon Cost
This cost (Cc,$) is simply the product of the initial carbon requirement (Mc)
and the current price of carbon. As adsorber vendors buy carbon in very
large quantities (million-pound lots or larger), their cost is somewhat lower
than the list price. A typical vendor cost is $2.00/lb (fall 1989 dollars).[8]
Thus:
Ce = 2.00MC (4.12)
4.3.1.2 Vessel Cost
The cost of an adsorber vessel is primarily determined by its dimensions
which, in turn, depend upon the amount of carbon it must hold and the
superficial gas velocity through the bed that must be maintained for opti-
mum adsorption. The desired superficial velocity is used to calculate, the
cross-sectional area of the bed perpendicular to the gas flow. An acceptable
superficial velocity is established empirically, considering desired removal
efficiency, the carbon particle size and bed porosity, and other factors. Fpr
example, one adsorber vendor recommends a superficial bed velocity of 85
ft/min[7], while an activated carbon manufacturer cautions against exceed-
ing 60 ft/min in systems operating at one atmosphere.[5] Another vendor
uses a 65 ft/min superficial face velocity in sizing its adsorber vessels.[8]
Lastly, there are practical limits to vessel dimensions which also influence
their sizing. That is, due to shipping restrictions, vessel diameters rarely
exceed 12 feet, while their length is generally limited to 50 feet.[8]
The cost of a vessel is usually correlated with its weight. However, as the
weight is often difficult to obtain or calculate, the cost may be estimated
from the external surface area. This is true because the vessel material
cost—and the cost of fabricating that material—is directly proportional
to its surface area. The surface area (S, ft2) of a vessel is a function of
its length (L, ft) and diameter (D, ft), which in turn, depend upon the
superficial bed face velocity, the L/D ratio, and other factors.
Most commonly, adsorber vessels are cylindrical in shape and erected
horizontally (as in Figure 4.1). Vessels configured in this manner are gen-
4-21
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erally subjected to the constraint that the carbon volume occupies no more
than 1/3 of the vessel volume [7,8]. It can be shown that this constraint
limits the bed depth to no more than
Maximum bed depth w - . (4-13)
\.£t
The vessel length, L, and diameter, D, can be estimated by solving two
relationships, namely, (1) the equation relating carbon volume, and thus
vessel volume, to L and D, and (2) the equation relating volumetric flow
rate, superficial velocity, and cross-section normal to flow. If one assumes
that the carbon bulk density is 30 lb/ft3, then one can show that:
z, - (4.14)
where
D = vessel diameter (ft)
L = vessel length (ft)
Vfc = bed superficial velocity (ft/min)
M'c = carbon requirement per vessel (Ibs)
Q' = volumetric flow rate per adsorbing vessel (acfm)
Because the constants in equations 4.14 and 4.15 are not dimensionless, one
must be careful to use the units specified in these equations.
Although other design considerations can result in different values of
L and D, these equations result in L and D which are acceptable from
the standpoint of "study" cost estimation for horizontal, cylindrical vessels
which are larger than 2-3 feet in diameter.
The carbon requirement and flow rate for each adsorber vessel can be
calculated as follows:
Mj = M<
(NA + ND)
Q
At gas flow rates (Q') of less than 9,000 scfm, it is usually more feasi-
ble to erect the adsorber vessels vertically instead of horizontally. [8] If so,
4-22
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the vessel diameter can be calculated from the volumetric flow rate per
adsorbing vessel and the bed superficial velocity as follows:
(4.16)
The vertical vessel length will depend principally on the carbon bed thick-
ness. Additional space must be included below the carbon bed for bed
support and above and below the bed for distribution and disengaging of
the gas stream and for physical access to the carbon bed. In smaller diame-
ter vessels, access to both sides of the bed is usually not required. However,
1 to l| feet must be provided on each side for gas distribution and disen-
gagement, or 2 to 3 feet overall. For longer vessels, 2 to 3 feet at each end
of the vessel is typically provided for access space.
Given the mass of carbon in the bed, the carbon bulk density, and the
bed diameter (i.e., the cross-sectional area normal to flow), determining the
carbon bed thickness is straightforward using the following equation:
__ volume of carbon __ .,
cross-sectional area normal to flow ' '
where
p = carbon bulk density (lb/ft3, assume 30 lb/ft3)
The vessel length is, therefore,
L = tb + ta>a (4.18)
where
<0lff = access/ gas distribution allowance
= 2 to 6 feet (depending on vertical vessel diameter)
Finally, xise the following equation to calculate the surface area of either
a horizontal or vertical vessel:
5 = nD(L + D/2) (4.19)
Similar equations can be developed for other vessel shapes, configurations,
etc.
Based on vendor data, we developed a correlation between adsorber
vessel cost and surface area: [8]
Cv = 271 5°-r78 (4.20)
4-23
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where C, = vessel cost (fall 1989 $), F.O.B. vendor
and 97 < S < 2,110ft2.
These units would be made of 304 stainless steel, which is the most
common material used in fabricating adsorber vessels.[7,8] However, to ob-
tain the cost of a vessel fabricated of another material, multiply Cv by an
adjustment factor (Fm). A few of these factors are listed below:
Material
Stainless steel, 316
Carpenter 20 CB-3
Monel-400
Nickel-200
Titanium
Fm Factor
1.3
1.9
2.3
3.2
4.5
Reference(s)
[7,8,9]
[9]
[7,9]
[9]
[9]
4.3.1.3 Total Purchased Cost
As stated earlier, the costs of such items as the fans, pumps, condenser,
decanter, instrumentation, and internal piping can be factored from the sum
of the costs for the carbon and vessels. Based on four data points derived
from costs supplied by an equipment vendor [8], we found that, depending
on the total gas flow rate (Q), the ratio (Rc) of the total adsorber equipment
cost to the cost of the vessels and carbon ranged from 1.14 to 2.24. These
data points spanned a gas flow rate range of approximately 4,000 to 500,000
acfm. The following regression formula fit these four points:
Re = 5.82 Q-°'133 (4.21)
where
4,000 < Q (acfm) < 500,000
Correlation coefficient (r) = 0.872
The total adsorber equipment cost (C^) would be the product of R,. and
the sum of the carbon and vessel costs, or:
CA = Rc [Ce + (NA + ND)CV] (4.22)
4-24
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4.3.1.4 Total Capital Investment
As discussed in Chapter 2, in the methodology used in this Manual, the
total capital investment (TCI) is estimated from the total purchased cost
via an overall direct/indirect installation cost factor. A breakdown of that
factor for carbon adsorbers is shown in Table 4.2. As Chapter 2 indicates,
the TCI also includes costs for land, working capital, and off-site facilities,
which are not included in the direct/indirect installation factor. However,
as these items are rarely required with adsorber systems, they will not be
considered here. Further, no factors have been provided for site preparation
(SP) and buildings (Bldg.), as these site-specific costs depend very little on
the purchased equipment cost.
Note that the installation factor is applied to the total purchased equip-
ment cost, which includes the cost of such auxiliary equipment as the stack
and external ductwork and such costs as freight and sales taxes (if applica-
ble). ("External ductwork" is that ducting needed to convey the exhaust
gas from the source to the adsorber system, and then from the adsorber
to the stack. Costs for ductwork and stacks are shown elsewhere in this
Manual.) Normally, the adjustment would also cover the instrumentation
cost, but this cost is usually included with the adsorber equipment cost.
Finally, note that these factors reflect "average" installation conditions and
could vary considerably, depending upon the installation circumstances.
4.3.2 Cannister Systems
Once the carbon requirement is estimated using the above procedure, the
number of cannisters is determined. This is done simply by dividing the
total carbon requirement (Me) by the amount of carbon contained by each
cannister (typically, 150 Ibs.). This quotient, rounded to the next high-
est digit, yields the required number of cannisters to control the vent in
question.
Costs for a typical cannister (Calgon's Ventsorb®) are listed in Ta-
ble 4.3. These costs include the vessel, carbon, and connections, but do
not include taxes, freight, or installation charges. Note that the cost per
unit decreases as the quantity purchased increases. Each cannister contains
Calgon's "BPL" carbon (4 x 10 mesh), which is commonly used in indus-
4-25
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Table 4.2: Capital Cost Factors for Carbon Adsorbers"
Cost Item Factor
Direct Costs
Purchased equipment costs
Adsorber + auxiliary equipment* As estimated, A
Instrumentation0 0.10 A
Sales taxes 0.03 A
Freight 0.05_A
Purchased equipment cost, PEC B = OcTA
Direct installation costs
Foundations & supports 0.08 B
Handling & erection 0.14 B
Electrical 0.04 B
Piping 0.02 B
Insulation 0.01 B
Painting 0.01 B
Direct installation costs 0.30 B
Site preparation As required, SP
Buildings As required, Bldg.
Total Direct Costs, DC 1.30 B + SP + Bldg.
Indirect Costs (installation)
Engineering 0.10 B
Construction and field expenses 0.05 B
Contractor fees 0.10 B
Start-up 0.02 B
Performance test 0.01 B
Contingencies 0.03 B
Total Indirect Costs, 1C 0.31 B
Total Capital Investment = DC + 1C 1.61 B + SP + Bldg.
"Reference [10].
Ductwork and any other equipment normally not included with unit furnished by
adsorber vendor.
instrumentation and controls often furnished with the adsorber, and thus included
in the EC.
4-26
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Table 4.3: Equipment Costs (Spring 1986 $) for a Typical Cannister
Adsorber"
Quantity Equipment Cost (each)6
1-3
4-9
10-29
>30
$687
659
622
579
b These costs are F.O.B., Pittsburgh,
PA. They do no* include taxes and
freight charges.
trial adsorption. However, to treat certain VOCs, more expensive speciality
carbons (e.g., "FCA 4 x 10") are needed. These carbons can increase the
equipment cost by 60% or more.[4] As is indicated in the caption of Table
4.3, these prices are in Spring 1986 dollars. Since then, however, the prices
of these cannisters have increased modestly—approximately 10%.[11]
As fewer installation materials and labor are required to install a can-
nister unit than a fixed-bed system, the composite installation factor is
consequently lower. The only costs required are those needed to place
the cannisters at, and connect them to, the source. This involves a small
amount of piping only; little or no electrical work, painting, foundations,
or-the like would be needed. Twenty percent of the sum of the czinnister(s)
cost, freight charges, and applicable sales taxes would cover this installation
cost.
4.4 Estimating Total Annual Cost
As Chapter 2 of this Manual explains, the total annual cost is comprised of
three components: direct coats, indirect costs, and recovery credits. These
will be considered separately.
4-27
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4.4.1 Direct Annual Costs
These include the following expenditures: steam, cooling water, electricity,
carbon replacement, operating and supervisory labor, and maintenance la-
bor and materials. Of these, only electricity and solid waste disposal would
apply to the cannister-type adsorbers.
4.4.1.1 Steam
As explained in section 4.1, steam is used during the desorption cycle. The
quantity of steam required will depend on the amount of carbon in the
vessel, the vessel dimensions, the type and amount of VOC adsorbed, and
other variables. Experience has shown that the steam requirement ranges
from approximately 3 to 4 Ibs of steam/lb of adsorbed VOC.[7,8] Using
the midpoint of this range, we can develop the following expression for the
annual steam cost:
C, = 3.50 x 10~3mvoc0, P, (4.23)
where
C, = steam cost ($/yr)
8, = system operating hours (h/yr)
mVOc = VOC inlet loading (Ibs/h)
p, = steam "price ($/thous. Ibs)
If steam price data are unavailable, one can estimate its cost at 120% of
the fuel cost. For example, if the local price of natural gas were $5.00/mil-
lion BTU, the estimated steam price would be $6.00/million BTU which is
approximately $6.00/thousand Ibs. (The 20% factor covers the capital and
annual costs of producing the steam.)
4.4.1.2 Cooling Water
Cooling water is consumed by the condenser in which the steam-VOC mix-
ture leaving the desorbed carbon bed is totally condensed. Most of the
condenser duty is comprised of the latent heat of vaporization (AH0) of the
steam and VOC. As the VOC AHW are usually small compared to the steam
AH, (about 1000 BTU/lb), the VOC AH, may be ignored. So may the
4-28
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sensible heat of cooling the water-VOC condensate from the condenser inlet
temperature (about 212°F) to the outlet temperature. Therefore, the cool-
ing water requirement is essentially a function of the steam usage and the
allowable temperature rise in the coolant, which is typically 30° to 40°F.[7]
Using the average temperature rise (35°F), we can write:
n
Cw = 3.43—PC* (4.24)
p*
where
Cow = cooling water cost ($/yr)
Pew — cooling water price ($/thous. gal.)
If the cooling water price is unavailable, use $0.15 to $0.30/thousand gal-
lons.
4.4.1.3 Electricity
In fixed-bed adsorbers, electricity is consumed by the system fan, bed dry-
ing/cooling fan, cooling water pump, and solvent pump(s). Both the system
and bed fans must be sized to overcome the pressure drop through the car-
bon beds. But, while the system fan must continuously convey the total
gas flow "through the system, the bed cooling fan is only used during a part
of the desorption cycle (one-half hour or less).
For both fans, the horsepower needed depends both on the gas flow and
the pressure drop through the carbon bed. The pressure drop through the
bed (APfc) depends on several variables, such as the adsorption tempera-
ture, bed velocity, bed characteristics (e.g., void fraction), and thickness.
But, for a given temperature and carbon, the pressure drop per unit thick-
ness depends solely on the gas velocity. For instance, for Calgon's "PCB"
carbon (4 x 10 mesh), the following relationship holds: [5]
= 0.03679UJ, + 1.107 x lO'X2 (4.25)
where
= pressure drop through bed (inches of water/foot
of carbon)
= superficial bed velocity (ft/min)
4-29
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As Equation 4.17 shows, the bed thickness (t&, ft) is the quotient of
the bed volume (V&) and the bed cross- sectional area (Aj,). For a 30 lb/ft3
carbon bed density, this becomes:
0.0333A/;
(For vertically erected vessels, A\, = Q'/v&> while for horizontally erected
cylindrical vessels, A «LD.) Once AP^ is known, the system fan horsepower
requirement (hpsf) can be calculated:
hpgf = 2.50 x 10~4 Q AP, (4.27)
where
Q = gas volumetric flow through system (acfm)
AP, = total system pressure drop = AP& 4- 1
(The extra inch accounts for miscellaneous pressure losses through the ex-
ternal ductwork and other parts of the system. [7] However, if extra long
duct runs and/or preconditioning equipment are needed, the miscellaneous
losses could be much higher.)
This equation incorporates a fan efficiency of 70% and a motor efficiency
of 90%, or 63% overall.
The horsepower requirement for the bed drying/cooling fan (hpcr) is
computed similarly. While the bed fan pressure drop would still be APj,,
the gas flow and operating times would be different. For typical adsorber
operating conditions, the drying/cooling air requirement would be 50 to 150
ft3/lb carbon, depending on the bed moisture content, required temperature
drop, and other factors. The operating time (0cf) would be the product of
the drying/cooling time per desorption cycle and the number of cycles per
year. It can be shown that:
0cf = OA9D(NAe./9A) (4.28)
(The "0.4" allows for the fact that as a rule-of- thumb, approximately 40%
of the desorption cycle is used for bed drying/cooling.)
The cooling water pump horsepower requirement (hpCWp) would be
computed as follows:
, 2.52 x 10-4gcw-ff s
- — - (4.29)
4-30
-------
where
gcw = cooling water flow (gal/min)
H = required head (nominally 100 feet of water)
s = specific gravity of fluid relative to water at 60°F
77 = combined pump-motor efficiency.
The annual operating hours for the cooling water pump (0Cwp) would
be computed using Equation 4.28, after substituting "0.6" for 0.4. The 0.6
factor accounts for the fact that the cooling water pump is only used during
the steaming portion of the regeneration cycle, while the condenser is in
operation.
Equation 4.29 may also be used to compute the solvent pump horse-
power requirement. In the latter case, the flow (q,) would be different, of
course, although the same head — 100 ft. of water — could be used. The spe-
cific gravity would depend on the composition and temperature of the con-
densed solvent. For example, the specific gravity of toluene at 100°F would
be approximately 0.86 at 70°F. (However, the solvent pump horsepower is
usually very small — usually < 0.1 hp. — so its electricity consumption can
usually be neglected.)
Once the various horsepowers are calculated, the electricity usage (in
kWh) is calculated, by multiplying each horsepower value by 0.746 (the
factor for converting hp to kilowatts) and the number of hours each fan
or pump operates annually. For the system fan, the hours would be the
annual operating hours for the system (0,). But, as discussed above, the
operating times for the bed drying/cooling fan and cooling water pump
would be different.
To obtain the annual electricity cost, simply multiply kWh'by the elec-
tricity price (in $/kWh) that applies to the facility being controlled.
For cannister units, use equation 4.27 to calculate the fan horsepower
requirement. However, instead of AP&, use the following to compute the
total cannister pressure drop ( APC, inches of water):[4]
= 0.047K?C + 9.29 x W~4Ql (4.30)
where Qc = flow through the cannister (acfrn).
4-31
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4.4.1.4 Carbon Replacement
As discussed above, the carbon has a different economic life than the rest
of the adsorber system. Therefore, its replacement cost must be calculated
separately. Employing the procedure detailed in Chapter 2, we have:
CRCe = CRFC(1MCC + Cd) (4-31)
where
CRFC = capital recovery factor for the carbon
1.08 = taxes and freight factor
Cc, Cci = initial cost of carbon (F.O.B. vendor) and carbon
replacement labor cost, respectively ($)
The replacement labor cost covers the labor cost for removing spent
carbon from vessels and replacing it with virgin or regenerated carbon.
The cost would vary with the amount of carbon being replaced, the labor
rates, and other factors. For example, to remove and replace a 50,000 pound
carbon charge would require about 16 person-days, which, at typical wage
rates, is equivalent to approximately $0.05/lb replaced. [12]
A typical life for the carbon is five years. However, if the inlet contains
VOCs that are very difficult to desorb, tend to polymerize, or react with
other constituents, a shorter carbon lifetime — perhaps as low as two years —
would be likely.fl] For a five-year life and 10% interest rate, CRFC = 0.2638.
4.4.1.5 Solid Waste disposal
Disposal costs are rarely incurred with fixed-bed adsorbers, because the
carbon is almost always regenerated in place, not discarded. In certain
cases, the carbon in cannister units is also regenerated, either off-site or at
a central regeneration facility on-site. However, most cannister adsorbers
are disposed of once they become saturated. The entire cannister—carbon,
drum, connections, etc.—is shipped to a secure landfill. The cost of landfill
disposal could vary considerably, depending on the number of cannisters
disposed of, the location of the landfill, etc. Based on data obtained from
two large landfills, for instance, the disposal cost would range from approx-
imately $35 to $65 per cannister excluding transportation costs.[13,14]
4-32
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4.4.1.6 Operating and Supervisory Labor
The operating labor for adsorbers is relatively low, as most systems are
automated and require little attention. One-half operator hour per shift
is typical.[10] The annual labor cost would then be the product of this
labor requirement and the operating labor wage rate ($/h) which, naturally,
would vary according to the facility location, type of industry, etc. Add to
this 15% to cover supervisory labor, as Chapter 2 suggests.
4.4.1.7 Maintenance Labor and Materials
Use 0.5 hours/shift for maintenance labor [10] and the applicable mainte-
nance wage rate. If the latter data are unavailable, estimate the mainte-
nance wage rate at 110% of the operating labor rate, as Chapter 2 suggests.
Finally, for maintenance materials, add an amount equal to the mainte-
nance labor, also per Chapter 2.
4.4.2 Indirect Annual Costs
These include such costs as capital recovery, property taxes, insurance,
overhead, and administrative costs ("G&A"). The capital recovery cost
is based on the equipment lifetime and the annual interest rate employed.
(See Chapter 2 for a thorough discussion of the capital recovery cost and the
variables that determine it.) For adsorbers, the system lifetime is typically
ten years, except for the carbon, which, as stated above, typically needs to
be replaced after five years. Therefore, when figuring the system capital
recovery cost, one should base it on the installed capital cost leas the coat of
replacing the carbon (i.e., the carbon cost plus the cost of labor necessary
to replace it). Substituting the initial carbon and replacement labor costs
from equation 4.31, we obtain:
CRC. = [TCI - (1.08Ce 4- Cel}} CRF. (4.32)
4-33
-------
where
CRC, = capital recovery cost for adsorber system ($/yr)
TCI = total capital investment ($)
1.08 = taxes and freight factor
CcjCd = initial carbon cost (F.O.B. vendor) and carbon
replacement cost, respectively ($)
CRF, = capital recovery factor for adsorber system (de-
fined in Chapter 2).
For a ten-year life and a 10% annual interest rate, the CRF, would be
0.1628.
As Chapter 2 indicates, the suggested factor to use for property taxes,
insurance, and administrative charges is 4% of the TCI. Finally, the over-
head is calculated as 60% of the sum of operating, supervisory, and main-
tenance labor, and maintenance materials.
The above procedure applies to cannister units as well, except that, in
most cases, the carbon is not replaced—the entire unit is. Cannisters are
generally used in specialized applications. The piping and ducting cost can
usually be considered a capital investment with a useful life of ten years.
However, whether the cannister itself would be treated as a capital or an
operating expense would'depend on the particular application and would
need to be evaluated on a case-by-case basis.
4.4.3 Recovery Credits
These apply to the VOC which is adsorbed, then desorbed, condensed, and
separated from the steam condensate. If the recovered VOC is sufficiently
pure, it can be sold. However, if the VOC layer contains impurities or
is a mixture of compounds, it would require further treatment, such as
distillation. Purification and separation costs are beyond the scope of this
chapter. Needless to say, the costs of these operations would offset the
revenues generated by the sale of the VOC. Finally, as an alternative to
reselling it, the VOC could be burned as fuel and valued accordingly. In
any case, the following equation can be used to calculate these credits:
RC = mvoc0,pvocE (4.33)
4-34
-------
where
RC = recovery credit ($/yr)
mvoc = VOC inlet loading (Ibs/h)
fft = system operating hours (h/yr)
pvoc = resale value of the recovered VOC ($/lb)
E = adsorber VOC control efficiency
By definition, the efficiency (E) is the difference between the inlet and
outlet VOC mass loadings, divided by the inlet loading. However, during an
adsorption cycle the outlet VOC loading will increase from essentially zero
at the start of the cycle to the breakthrough concentration at the end of the
cycle. Because the efficiency is a function of time, it should be calculated
via integration over the length of the adsorption cycle. To do this would
require knowledge of the temporal variation of the outlet loading during
the adsorption cycle. If this knowledge is not available to the Manual user,
a conservative approximation of the efficiency may be made by setting the
outlet loading equal to the breakthrough concentration.
4.4.4 Total Annual Cost
Finally, as explained in Chapter 2, the total annual cost (TAG) is the sum
of the' direct and indirect annual costs, less any recovery credits, or:
TAG = DC + IC-RC (4.34)
4.4.5 Example Problem
A source at a printing plant emitting 100 Ib/h of toluene is to be controlled
by a carbon adsorber. The plant proposes to operate the adsorber in a
continuous mode for 8,640 h/yr (360 days). While operating, two carbon
beds will be adsorbing, while a third will be desorbing/on stand by. For its
convenience, the plant has selected adsorption and desorption times of 12
and 5 hours, respectively. The total waste gas flow is 10,000 acfm at the
adsorber inlet conditions (one atmosphere and 77°F). The waste gas con-
tains negligible quantities of particulate matter and moisture. Further, the
applicable VOC regulation requires the adsorber to achieve a mean removal
efficiency of 98% during the entire adsorption cycle. Finally, assume that
4-35
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the recovered toluene is recycled at the source. Estimate the total capital
investment and total annual cost for the adsorber system.
Carbon Working Capacity: At the stated flow and pollutant load-
ing, the toluene inlet concentration is 710 ppm. This corresponds to a
partial pressure of 0.0104 psia. Substituting this partial pressure and the
toluene isotherm parameters (from Table 4.1) into equation 4.1, we obtain
an equilibrium capacity of 0.333 Ib/lb. By applying the rule-of-thumb dis-
cussed above (page 4-19), we obtain a working capacity of 0.167 Ib/lb (i.e.,
0.333/2).
Carbon Requirement: As stated above, this adsorber would have two
beds on-line and a third off-line. Is this a reasonable assumption? Equation
4.8 can answer this question. Substitution of the adsorption time and
numbers of adsorbing and desorbing beds yields:
Desorption time = 9D < 9A(ND/NA} = 12 h(l/2) = 6 h.
Because, the stated desorption time (5 hours) is less than 6 hours, the pro-
posed bed configuration is feasible. Next, calculate the carbon requirement
(Mc) from equation 4.10:
From equation 4.12, the carbon cost is:
Adsorber Vessel Dimensions and Cost: Assume that the vessels will
be erected horizontally and select a superficial bed velocity (vj,) of 75 ft/min.
Next, calculate the vessel diameter (D), length (L), and surface area (S)
from equations 4.14, 4.15, and 4.19, respectively. [Note: In these equations,
M'e = Me/(NA + ND) = 3,600 Ib and Q' = Q/NA = 5,000 acfm.]
= 0.127(3, 600)(75) = Q ^ ft
7.87 Q'\* 7.87
3,600 75
4-36
-------
5 = nD(L + D/2) = 283 ft2
Because S falls between 97 and 2,110 ft2, equation 4.20 can be used to
calculate the cost per vessel, Cv (assuming 304 stainless steel construction).
Thus:
Adsorber Equipment Cost: Recall that the adsorber equipment cost
is comprised of the adsorber vessels, carbon, and the condenser, decanter,
fan, pumps and other equipment usually included in the adsorber price.
The cost of the latter items are "factored" from the combined cost of the
vessels and carbon. Combining equations 4.21 and 4.22, we have:
CA = 5.82
-------
Table 4.4: Capital Costs for Carbon Adsorber System
Example Problem
Cost Item Cost
Direct Costs
Purchased equipment costs
Adsorber vessels and carbon $149,300
Auxiliary equipment 32,200
Sum = A $181,500
Instrumentation, 0.1 A° —
Sales taxes, 0.03A 5,450
Freight, 0.05A 9,080
Purchased equipment cost, B $196,000
Direct installation costs
Foundation and supports, 0.08B 15,680
Handling & erection, 0.14B 27,440
Electrical, 0.04B 7,840
Piping, 0.02B 3,920
Insulation for ductwork, 0.01B 1,960
Painting, 0.01B 1,960
Direct installation cost $58,800
Site preparation —
Facilities and buildings —
Total Direct Cost $254,800
Indirect Costs (installation)
Engineering, 0.10B 19,600
Construction and field expenses, 0.05B 9,800
Contractor fees, 0.1 OB 19,600
Start-up, 0.02B 3,920
Performance test, 0.01B 1,960
Contingencies, 0.03B 5,880
Total Indirect Cost $60,760
Total Capital Investment (rounded) $316,000
"The cost for this is included in the adsorber equipment cost.
4-38
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And:
Total Capital Investment (rounded) = 1.61 x MB" = $316,000
Annual Costs: Table 4.5 gives the direct and indirect annual costs for
the carbon adsorber system, as calculated from the factors in Section 4.4.
Except for electricity, the calculations in the table show how these costs
were derived. The following discussion will deal with the electricity cost.
First, recall that the electricity includes the power for the system fan,
bed drying/cooling fan, and the cooling water pump. (The solvent pump
motor is normally so small that its power consumption may be neglected.)
These consumptions are calculated as follows:
• System fan: From equation 4.27:
kWhgf = 0.746kW/hp x 2.50 x 10~4QAP, x 0,
But:
AP, (inches water) = APb + 1 = <6(0.03679vfc + 1.107 x 10~X) + 1
(The latter expression was derived from equation 4.25, assuming that
the carbon used in this example system is Calgon's "PCB", 4 x 10
mesh size.)
By assuming a carbon bed density of 30 lb/ft3, Equation 4.26 can be
used to calculate the bed thickness (
-------
Table 4.5: Annual Costs for Carbon Adsorber System
Example Problem
Cost Item
Calculations
Cost
Direct^ Annual Costs, DC
Operating Labor
Operator
Supervisor
Operating materials
Maintenance
Labor
Material
0.
360 days
$12
~h~
15% of operator = .15 x 6,480
0.5 h 3sh x 360 days
sKIff x clay x yr
$13.20
n
100% of maintenance labor
Replacement parts, carbon (5 year life)
Replacement labor 0.2638 ($0.05/lb x 10,800 Ib)
0.2638($21,600x 1.08)
8640 h
$0.06 /kWh x 131,000 kWh/yr
3.5 Ib
Ib VOC
3.43 gal
(3.5 x 100 x 8640) Ib steam $020
Ib steam yr x IQ* gai
$6 „ 100 Ib VQC
x l(FIb H
Carbon cost0
Utilities
Electricity
Steam
Cooling water
Total DC
Indirect Annual Costs, 1C
Overhead 60% of sum of operating, supv., k maint. labor
k maint. materials = 0.6(6,480 + 970 + 7,130 +
7,130)
Administrative charges 2% of Total Capital Investment = 0.02($316,000)
1% of Total Capital Investment = 0.01($316,000)
1% of Total Capital Investment = 0.01($316,000)
0.1628(316,000 - 0.05(10,800) - 1.08(21,600)]
Property tax
Insurance
Capital recovery*
Total 1C
Recovery Credit (toluene) .98($°-055-3
Total Annual Cost (rounded)
yf
86,480
970
7,130
7,130
140
6,150
7,860
18,140
2,070
$56,070
13,030
6,320
3,160
3,160
47,560
$73,230
(46,820)
$82,500
0 The 1.08 factor is for freight and sales taxes.
* The capital recovery cost factor, CRF, is a function of the adsorber or equipment life and
the opportunity cost of the capital (i.e., interest rate). For example, for a 10 year equipment
life and a 10% interest rate, CRF = 0.1628.
4-40
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• Bed drying/cooling fan: During the drying/cooling cycle, the pressure
drop through the bed also equals AP&. However, as section 4.4.1.3
indicates, the flow and operating time are different. For the air flow,
take the midpoint of the range given on page 4-30 (100 ft3 air/lb
carbon) and divide by 2 hours (the bed drying/cooling time), yielding:
100 ft3/lb x 3,600 Ibs x 1/120 min = 3,000 acfm. Substituting this
into equation 4.27 results in:
2.50 x 10~4 x 7.09 inches x 3,000 acfm = 5.32 hp
From equation 4.28, we get:
0cf = (0.4)(5 h)(2)(8,640 h)/12 h = 2,880 h
Thus:
kWhcf = 0.746 kW/hp x 5.32 hp x 2,880 h = 11,400 kWh/yr
• Cooling water pump: The cooling water pump horsepower is calcu-
lated from equation 4.29. Here, let 77 = 63% and H = 100 ft. The
cooling water flow (qcw) is the quotient of the annual cooling water
requirement and the annual pump operating time. From the data
in Table 4.5, we obtain the cooling water requirement: 10,400,000
gal/yr. The pump annual operating time is obtained from equation
4.28 (substituting 0.6 for 0.4), or 0Cwp = (0.6)(5 h)(2)(8,640)/12 =
4,320 h/yr.
Thus:
(2.52 x 10-4)(100ft) 10,400, OOP gal/yr _, flnl.
= 0^3 X 4,320h/yrx60min/yr " 1>6°hp
And:
kWhcwp = 0.746 kW/h x 1.60 hp x 4,320 h/yr = 5,160 kWh/yr
Summing the individual power consumptions, we get the value shown
in Table 4.5: 131,000 kWh/yr.
Recovery Credit: As Table 4.5 indicates, a credit for the recovered
toluene has been taken. However, to account for miscellaneous losses and
contamination, the toluene is arbitrarily valued at one-half the current
(November 1989) market price ($0.0533/lb = $lll/ton).[15]
4-41
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Total Annual Cost: The sum of the direct and indirect annual costs,
less the toluene recovery credit, yields a net total annual cost of $82,500.
Clearly, this "bottom line" is very sensitive to the recovery credit and, in
turn, the value given the recovered toluene. For instance, if it had been
valued at the full market price ($221/ton), the credit would have doubled
and the total annual cost would have been a net credit of $35,700. Thus
when incorporating recovery credits, it is imperative to select the value of
the recovered product carefully.
4-42
-------
-------
References
[1] Correspondence: Robert L. Stallings and William Klotz (Research Tri-
angle Institute, Research Triangle Park, NC) to William M. Vatavuk
(U. S. EPA, OAQPS, Research Triangle Park, NC), June 24, 1986.
[2] Calvert, Seymour and Englund, Harold M. (eds.), Handbook of Air
Pollution Control Technology, John Wiley & Sons, New York, 1984,
pp. 135-192.
[3] Handbook of Chemistry and Physics, 54th Edition, The Chemical Rub-
ber Company, Cleveland, 1973-74, pp. D85-D92.
[4] "Calgon Ventsorb® for Industrial Air Purification" (Bulletin 23-56a),
Calgon Corporation, Pittsburgh, 1986.
[5] Adsorption Handbook, Calgon Corporation, Pittsburgh, 1980.
[6] Rogers, Tony, "Comparison of BED-SIZE and Calgon Adsorption
Isotherms", Research Triangle Institute (Research Triangle Park, NC),
January 20, 1988.
[7] Correspondence: Richard Selznick (Baron Blakeslee, Inc., Westfield,
NJ) to William M. Vatavuk (U. S. EPA, OAQPS, Research Triangle
Park, NC), April 23, 1986.
[8] Correspondence: Denny Clodfelter (M&W Industries, Inc., Rural Hall,
NC) to William M. Vatavuk (U. S. EPA, OAQPS, Research Triangle
Park, NC), September 25, 1989.
[9] Matley, Jay (ed.), Modern Cost Engineering, McGraw-Hill Publica-
tions Co., New York, 1984, p. 142.
4-43
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[10] Vatavuk, William M. and Neveril, Robert, "Estimating Costs of Air
Pollution Control Systems, Part II: Factors for Estimating Capital
and Operating Costs," Chemical Engineering, November 3, 1980, pp.
157-162.
[11] Telephone conversation: Robert Bradley (Calgon Corporation, Char-
lotte, NC) with William M. Vatavuk (U. S. EPA, OAQPS, Research
Triangle Park, NC), December 5, 1989.
[12] Telephone conversation: Robert L. S tailings (Research Triangle Insti-
tute, Research Triangle Park, NC) with William M. Vatavuk (U. S.
EPA, OAQPS, Research Triangle Park, NC), September 11, 1986.
[13] Correspondence: William Kitto (Chemwaste, Sulphur, LA) to William
M. Vatavuk (U. S. EPA, OAQPS, Research Triangle Park, NC), July
25, 1986.
[14] Correspondence: Jerry Locklear (GSX, Pinewood, SC) to William M.
Vatavuk (U. S. EPA, OAQPS, Research Triangle Park, NC), July 25,
1986.
[15] Chemical Marketing Reporter, December 2, 1989.
4-44
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-f
Chapter 5
FABRIC FILTERS
James H. Turner
Andrew S. Viner
Research Triangle Institute
Research Triangle Park, N.C. 22709
John D. McKenna
ETS, Inc.
Roanoke, VA 24018-4394
Richard E. Jenkins
William M. Vatavuk
Standards Development Branch, OAQPS
U.S. Environmental Protection Agency
Research Triangle Park, N.C. 22711
November, 1989
5-1
-------
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Contents
5.1 Process Description 5-4
5.1.1 Introduction 5-4
5.1.2 Types of Fabric Filters 5-5
5.1.2.1 Shaker Cleaning 5-5
5.1.2.2 Reverse-air Cleaning 5-6
5.1.2.3 Pulse-jet Cleaning 5-7
5.1.3 Auxiliary Equipment 5-8
5.1.4 Fabric Filtration Theory 5-9
5.1.4.1 Reverse Air/Shake Deflate Baghouses .... 5-10
5.1.4.2 Pulse-Jet Baghouses 5-13
5.2 Design Procedures 5-16
5.2.1 Gas-to-Cloth Ratio 5-16
5.2.1.1 Gas-to-Cloth Ratio From Similar Applications 5-16
5.2.1.2 Gas-to-Cloth Ratio From Manufacturer's Meth-
ods 5-19
5.2.1.3 Gas-to-Cloth Ratio From Theoretical/Empirical
Equations 5-19
5.2.2 Pressure Drop 5-23
5.2.3 Particle Characteristics 5-24
5.2.4 Gas Stream Characteristics 5-24
5.2.4.1 Temperature 5-25
5-2
-------
5.2.4.2 Pressure 5-25
5.2.5 Equipment Design Considerations 5-27
5.2.5.1 Pressure or Suction Housings 5-27
5.2.5.2 Standard or Custom Construction 5-27
5.2.5.3 Filter Media 5-28
5.3 Estimating Total Capital Investment 5-29
5.3.1 Equipment Cost 5-29
5.3.1.1 Bare Baghouse Costs 5-29
5.3.1.2 Bag Costs 5-35
5.3.1.3 Auxiliary Equipment 5-38
5.3.2 Total Purchased Cost 5-38
5.3.3 Total Capital Investment 5-38
5.4 Estimating Total Annual Costs 5-39
5.4.1 Direct Annual Cost 5-39
5.4.1.1 Operating and Supervisory Labor 5-41
5.4.1.2 Operating Materials 5-41
5.4.1.3 Maintenance 5-41
5.4.1.4 Replacement Parts 5-41
5.4.1.5 Electricity 5-42
5.4.1.6 Fuel 5-43
5.4.1.7 Water 5-43
5.4.1.8 Compressed Air 5-44
5-3
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5.4.1.9 Dust Disposal 5-44
5.4.2 Indirect Annual Cost 5-44
5.4.3 Recovery Credits 5-45
5.4.4 Total Annual Cost 5-45
5.4.5 Example Problem 5-45
5.5 Acknowledgment 5-51
References 5-52
5.1 Process Description
5.1.1 Introduction
A fabric filter unit consists of one or more isolated compartments containing
rows of fabric filter bags or tubes. Particle-laden gas passes up (usually)
along the surface of the bags then radially through the fabric. Particles are
retained on the upstream face of the bags, while the cleaned gas stream
is vented to the atmosphere. The filter is operated cyclically alternating
between relatively long periods of filtering and short periods of cleaning.
During cleaning, dust that has accumulated on the bags is removed from
the fabric surface and deposited in a hopper for subsequent disposal.
Fabric filters will collect particle sizes ranging from submicron to several
hundred microns in diameter at efficiencies generally in excess of 99 or 99.9
percent. The dust cake collected on the fabric is primarily responsible for
such high efficiency. Gas temperatures up to about 500°F, with surges to
about 550° F can be accommodated routinely. Most of the energy used to
operate the system appears as pressure drop across the bags and associated
hardware and ducting. Typical values of pressure drop range from about
5 to 20 inches of water. Fabric filters are used where high-efficiency par-
ticle collection is required. Limitations are imposed by gas characteristics
(temperature and corrosivity) and particle characteristics (primarily sticki-
ness) that affect the fabric or its operation'and that cannot be economically
5-4
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accommodated.
Important process variables include particle characteristics, gas charac-
teristics, and fabric properties. The most important design parameter is
the air- or gas-to-cloth ratio, and the usual operating parameter of interest
is pressure drop across the filter system. The major operating feature of
fabric filters that distinguishes them from other gas filters is the ability to
renew the filtering surface periodically by cleaning.
Another type of fabric filter currently being developed is the electro-
statically enhanced filter. Pilot plant baghouses employing this new tech-
nology have shown substantially lower pressure drops than conventional
filter designs. Further, some cost analyses have shown that electrostati-
cally enhanced baghouses could have lower lifetime costs than convention
baghouses. The purpose of this chapter, however, is to focus only on cur-
rently available commercial filters. Anyone interested in electrostatically
enhanced filtration may consult such references as Van Osdell et al [1],
Viner et al [2], or Donovan [3].
In this section, the types of fabric filters and the auxiliary equipment
required are discussed first from a general viewpoint. Then, fabric filtration
theory as applied to each type of filter is discussed to lay a foundation for
the sizing procedures outlined in the Section 5.2.
5.1.2 Types of Fabric Filters
Fabric filters can be categorized by several means, including type of clean-
ing (shaker, reverse-air, pulse-jet), direction of gas flow (from inside the
bag towards the outside or vice versa), location of the system fan (suction
or pressure), or size (low, medium, or high gas flow quantity). Of these
four approaches, the cleaning method is probably the most distinguishing
feature. Fabric filters are discussed in this section based on the type of
cleaning employed.
5.1.2.1 Shaker Cleaning
For any type of cleaning, enough energy must be imparted to the fabric to
overcome the adhesion forces holding dust to the bag. In shaker cleaning,
5-5
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used with inside to outside gas flow, this is accomplished by suspending the
bag from a motor-driven hook or framework that oscillates. Motion may
be imparted to the bag in several ways, but the general effect is to create a
sine wave along the fabric. As the fabric moves outward, accumulated dust
on the surface moves with the fabric. When the fabric reaches the limit
of its extension, the patches of dust have enough inertia to tear away from
the fabric and descend to the hopper.
For small, single-compartment baghouses, a lever attached to the shaker
mechanism may be operated manually at appropriate intervals, typically at
the end of a shift. In multi-compartment baghouses, a timer or a pressure
sensor responding to system pressure drop initiates bag shaking automat-
ically. The compartments operate in sequence so that one compartment
at a time is cleaned. Forward gas flow to the compartment is stopped,
dust is allowed to settle, residual gas flow stops, and the shaker mecha-
nism is switched on for several seconds to a minute or more. The settling
and shaking periods may be repeated, then the compartment is brought
back on-line for filtering. Many large-scale shaker systems employ a small
amount of reverse air during the shaker cycle to assist cleaning by deflating
the bags.
Parameters that affect cleaning include the amplitude and frequency
of the shaking motion and the tension of the mounted bag. The first two
parameters are part of the baghouse design and generally are not changed
easily. The tension is set when bags are installed. Typical values are about
4 Hz for frequency and 2 to 3 inches for amplitude (half-stroke).[4] Some
installations allow easy adjustment of bag tension, while others require that
the bag be loosened and reclamped to its attaching thimble.
The vigorous action of shaker systems tends to stress the bags and
requires heavier and more durable fabrics. In the United States, woven
fabrics are used almost exclusively for shaker cleaning. [5] European practice
allows the use of felted fabrics at somewhat higher filtering velocities.
5.1.2.2 Reverse-air Cleaning
When glass fiber fabrics were introduced, a gentler means of cleaning the
bags was needed to prevent premature degradation. Reverse-air cleaning
was developed as-a less intensive way to impart energy to the bags. In this
5-6
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method, gas flow to the bags is stopped in the compartment being cleaned,
and a reverse flow of air is directed through the bags. This reversal of gas
flow gently collapses the bags, and dust is removed from the fabric surface
by shear forces developed between the dust and fabric as the latter changes
its contours. Other differences between reverse-air and shaker cleaning are
the installation of metal caps as an integral part of the bag and sewn-in
rings to prevent complete collapse of the bag, which may be 30 feet long,
during cleaning. Without these rings, collected dust tends to choke the
bag as the fabric collapses in on itself during inside-to-outside filtering. As
with multi-compartment shaker baghouses, the same cycle takes place in
reverse-air baghouses of stopping forward gas flow and allowing dust to
settle before cleaning action begins.
The source of reverse air is generally a separate system fan capable of
supplying clean, dry air for one or two compartments at a gas-to-cloth ratio
similar to that of the forward gas flow.
5.1.2.3 Pulse-jet Cleaning
This form of cleaning uses compressed air to force a burst of air down
through the bag and expand it violently. As with shaker baghouses, the
fabric reaches its extension limit and the dust separates from the bag. In
pulse jets, however, filtering gas flows are opposite in direction when com-
pared with shaker or reverse-air baghouses. Bags are mounted on wire
cages to prevent collapse while the dusty gas flows from outside the bag
to the inside. Instead of attaching both ends of the bag to the baghouse
structure, the bag and cage assembly generally is attached only at the top.
The bottom end of the assembly tends to move in the turbulent gas flow
and may contact other bags, which accelerates wear.
Although some pulse-jet baghouses are compartmented, most are not.
Bags are cleaned one row at a time when a timer initiates the burst of
cleaning air through a quick-opening valve. A pipe across each row of bags
carries the compressed air. The pipe is pierced above each bag so that
cleaning air exits directly downward into the bag. Some systems direct the
air through a short venturi that is intended to entrain additional cleaning
air. The pulse opposes and interrupts forward gas flow only for a few tenths
of a second. However, the quick resumption of forward flow redeposits most
of the dust back on the clean bag or on adjacent bags. An advantage of
5-7
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Fabric Filttr
Diraet Exhaust
YY 1
Slack
Duit Ramoval
Maehankai Collactor
Figure 5.1: Typical alternative auxiliary equipment items used with fabric
filter control systems.
pulse-jet cleaning is the reduction in baghouse size allowed by not having
to build an extra compartment for off-line cleaning.
5.1.3 Auxiliary Equipment
The typical auxiliary equipment associated with fabric filter systems is
shown in Figure 5.1. Along with the fabric filter itself, a control system
typically includes the following auxiliary equipment: a capture device (i.e.,
hood or direct exhaust connection); ductwork; dust removal equipment
(screw conveyor, etc.); fans, motors, and starters; and a stack. In addition,
spray chambers, mechanical collectors, and dilution air ports may be needed
to precondition the gas before it reaches the fabric filter. Capture devices
are usually hoods that exhaust pollutants into the ductwork or direct ex-
haust couplings attached to a process vessel. Hoods are more common,
yet poorly designed hoods will allow pollutants to escape. Direct exhaust
couplings are less common, requiring sweep air to be drawn through the
process vessel, and may not be feasible in some processes. Ductwork pro-
vides a means of moving the exhaust stream to the control device. Spray
chambers and dilution air ports are used to decrease the temperature of the
5-8
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pollutant stream to protect the filter fabric from excessive temperatures.
When a substantial portion of the pollutant loading consists of relatively
large particles, mechanical collectors such as cyclones are used to reduce
the load on the fabric filter itself. The fans provide motive power for air
movement and can be mounted before (pressure baghouse) or after (suction
baghouse) the filter. A stack, when used, vents the cleaned stream to the
atmosphere. Screw conveyors are often used to remove captured dust from
the bottom of the hoppers. Air conveying (pneumatic) systems and direct
dumping into containers are also used.
5.1.4 Fabric Filtration Theory
The key to designing a baghouse is to determine the face velocity that
produces the optimum balance between pressure drop (operating cost) and
baghouse size (capital cost). Major factors that affect design face velocity
(or gas-to-cloth ratio), discussed in Section 5.2, include particle and fabric
characteristics and gas temperature.
Although collection efficiency is another important measure of baghouse
performance, it is generally assumed that a properly designed and well run
baghouse will be highly efficient. Therefore, the design process focuses on
the pressure drop.
There are several contributions to the pressure drop across a baghouse
compartment, including the pressure drop from the flow through the in-
let and outlet ducts, from flow through the hopper regions, and from flow
through the bags. The pressure drop through the baghouse compartment
(excluding the pressure drop across the bags) depends largely on the bag-
house design and ranges from 1 to 2 inches of HaO[3] in conventional designs
and up to about 3 inches of H20 in designs having complicated gas flow
paths. This loss can be kept to a minimum (i.e., 1 inch of I^O or less) by
investing in a flow modeling study of the proposed design. A study of this
sort would cost on the order of $50,000 (in 1986). The pressure drop across
the bags (also called the tube-sheet pressure drop) can be as high as 10
inches of HjO or more. The tube-sheet pressure drop is a complex function
of the physical properties of the dust and the fabric and the manner in
which the baghouse is designed and operated. The duct and hopper losses
are constant and can be minimized effectively through proper design based
5-9
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on a knowledge of the flow through the baghouse.1
Fabric filtration is inherently a batch process that has been adapted
to continuous operation through clever engineering. One requirement for a
continuously operating baghouse is that the dust collected on the bags must
be removed periodically. Shaker and reverse-air baghouses are similar in
the sense that they both normally use woven fabric bags, run at relatively
low face velocities, and the filtration mechanism is cake filtration. That
is, the fabric merely serves as a substrate for the formation of a dust cake
that is the actual filtration medium. Pulse-jet baghouses generally use felt
fabrics and run with a high gas-to-cloth ratio (about double that of shaker
or reverse-air baghouses). The felt fabric may play a much more active role
in the filtration process. This distinction between cake filtration and fabric
filtration has important implications for the rate of pressure loss across the
filter bags. The theoretical description of cake filtration is quite different
from that for fabric filtration, and the design processes are quite different.
Accurate fabric selection is aided by bench-scale filtration tests at less than
one-tenth the cost of flow modeling. These tests can be used to investigate
fabric effects on pressure drop, cake release during cleaning, and collec-
tion efficiency. Electrical properties of the fabric, such as resistivity and
triboelectric order, may be measured to aid in fabric selection. Although
their effects are generally poorly understood, electrical/electrostatic effects
influence cake porosity and particle adhesion.[6,7,8]
The general equations used to size a baghouse follow beginning with the
reverse air/shake deflate type of baghouse.
5.1.4.1 Reverse Air/Shake Deflate Baghouses
The construction of a baghouse begins with a set of specifications including
average pressure drop, total gas flow, and other requirements; a maxi-
mum pressure drop may also be specified. Given these specifications, the
designer must determine the maximum face velocity that can meet these
requirements. The standard way to relate baghouse pressure drop to face
velocity is given by the relation:
(5.1)
*A procedure for estimating duct pressure losses is given in the "Ductwork" chapter of
this Manual.
5-10
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where
AP(0) = the pressure drop across the filter, a function of time,
e (in. H20)
S«y«(0) = system drag, a function of time [in. H20/(ft/niin)]
Va»« = average (i.e., design) face velocity, constant (ft/min)
For a multi-compartment baghouse, the system drag, which accounts for
most of the drag from the inlet flange to the outlet flange of the baghouse,
is determined as a combination of resistances representative of several com-
partments. For the typical case where the pressure drop through each
compartment is the same, it can be shown that:[16]
M
JM
where
M i r1 i M
M = number of compartments in the baghouse
= drag across compartment i
The compartment drag is a function of the amount of dust collected on the
bags in that compartment. In general, the dust will be distributed in a very
nonuniform manner. That is, there will be a variation of dust load from
one bag to the next and within a given bag there will also be a variation of
dust load from one area to another. For a sufficiently small area j within
compartment i, it can be assumed that the drag is a linear function of dust
load:
(5.3)
where
Se = drag of a dust-free filter bag [in. H2O/(ft/min)]
K2 = dust cake flow resistance {[in. H20/(ft/min)]/(lb/ft2)}
Wt,j(0) = dust mass per unit area of area j in compartment t,
"areal density" (lb/ft2)
If there are N different areas of equal size within compartment t, each with
a different drag Sy, then the total drag for compartment i can be computed
in a manner analogous to Equation 5.2:
5-11
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(5-4)
The constants Se and K2 depend upon the fabric and the nature and size of
the dust. The relationships between these constants and the dust and fabric
properties are not understood well enough to permit accurate predictions
and so must be determined empirically, either from prior experience with
the dust/fabric combination or from laboratory measurements. The dust
mass as a function of time is defined as:
!' Cin ViM d9 (5.5)
•Jo
WiM = Wr + C
where
Wr = dust mass per unit area remaining on a "clean" bag (lb/ft2)
Cin = dust concentration in the inlet gas (lb/ft3)
= face velocity through area j of compartment » (ft/min)
It is assumed that the inlet dust concentration and the filter area are con-
stant. The face velocity (gas-to-cloth ratio) through each filter area j and
compartment i changes with time, starting at a maximum value just after
cleaning and steadily decreasing as dust builds up on the bags. The indi-
vidual compartment face velocities are related to the average face velocity
by the expression:
.
•"•*
(,R,
(5.6)
compartments with equal area)
Equations 5.1 through 5.6 reveal that there is no explicit relationship be-
tween the design face velocity and the tube-sheet pressure drop. On the
contrary, the pressure drop that results from a given design can only be de-
termined by the simultaneous solution of Equations 5.1 through 5.5, with
Equation 5.6 as a constraint on that solution. This conclusion has several
implications for the design process. The design requires an iterative proce-
dure: one must begin with a known target for the average pressure drop,
propose a baghouse design (number of compartments, length of filtration
5-12
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period, etc.), assume a face velocity that will yield that pressure drop, and
solve the system of Equations 5.1 through 5.6 to verify that the calculated
pressure drop equals the target pressure drop. This procedure is repeated
until the specified face velocity yields an average pressure drop (and max-
imum pressure drop, if applicable) that is sufficiently close to the design
specification.
5.1.4.2 Pulse-Jet Baghouses
The distinction between pulse-jet baghouses and reverse-air and shaker bag-
houses is basically the difference between cake filtration and composite
dust/fabric filtration (noncake filtration). This distinction is more a mat-
ter of convenience than physics. In reality, pulse-jet baghouses have been
designed to operate in a variety of modes. Some pulse jets remain on-line at
all times and are cleaned frequently. Others are taken off-line for cleaning
at relatively long intervals. Obviously, if a compartment remains on-line
long enough without being cleaned, then the filtration mechanism becomes
that of cake filtration. A complete model of pulse-jet filtration therefore
must account for the depth filtration occurring on a relatively clean pulse-
jet filter, the cake filtration that inevitably results from prolonged periods
on-line, and the transition period between the two regimes.
Besides the question of filtration mechanism, there is also the question
of cleaning method. If a compartment is taken off-line for cleaning, then
the dust that is removed from the bags will fall into the dust hopper before
forward gas flow resumes. If a compartment is cleaned while on-line, then
only a small fraction of the dust removed from the bag will fall to the
hopper. The remainder of the dislodged dust will be redeposited (i.e.,
"recycled") on the bag by the forward gas flow. The redeposited dust
layer has different pressure drop characteristics than the freshly deposited
dust. The modeling work that has been done to date focuses on the on-line
cleaning method. Dennis and Klemm[9] proposed the following model of
drag across a pulse-jet filter:
5 = Se + (K2)eWc + K2W0 (5.7)
5-13
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where
S = drag across the filter
Se = drag of a just-cleaned filter
(K2)c = specific dust resistance of the recycling dust
Wc = area! density of the recycling dust
K2 = specific dust resistance of the freshly deposited dust
W0 = areal density of the freshly deposited dust
This model has the advantage that it can easily account for all three regimes
of filtration in a pulse-jet baghouse. As in Equations 5.1 to 5.6, the drag,
filtration velocity and area! densities are functions of time, 9. For given
operating conditions, however, the values of Se, (K2)c, and Wc may be
assumed to be constant, so that they can be grouped together:
AP = (PE)*W + K,W0V (5.8)
where
AP = pressure drop (in. H20)
V = filtration velocity (ft/min)
(PE)Aw = [S. + (K2)CWC] V
Equation 5.8 describes the pressure drop bejiavior of an individual bag. To
extend this single bag result to a multiple-bag compartment, Equation 5.7
would be used to determine the individual bag drag and the total baghouse
drag would then be computed as the sum of the parallel resistances. Pres-
sure drop would then be calculated as in Equation 5.1. It seems reasonable
to extend this analysis to the case where the dust is distributed unevenly
on the bag and then apply Equation 5.7 to each area on the bag, followed
by an equation analogous to 5.4 to compute the overall bag drag. The diffi-
culty in doing this is that one must assume values for Wc for each different
area to be modeled.
The disadvantage of the model represented by Equations 5.7 and 5.8
is that the constants, Se, K2, and Wc, cannot be predicted at this time.
Consequently, correlations of laboratory data must be used to determine
the value of (PE)^. For the fabric-dust combination of Dacron felt and coal
fly ash, Dennis and Klemm[9] developed an empirical relationship between
(PE)^, the face velocity, and the cleaning pulse pressure. This relationship
(converted from metric to English units) is as follows:
5-14
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r0-88 (5.9)
where
V/ = face velocity (ft/min)
PJ = pressure of the cleaning pulse
(usually 60 to 100 psig; see Section 5.4.1)
This equation is essentially a regression fit to a limited amount of lab-
oratory data and should not be applied to other dust/fabric combinations.
The power law form of Equation 5.9 may not be valid for other dusts or fab-
rics. Consequently, more data should be collected and analyzed before the
model represented by Equation 5.9 can be used for rigorous sizing purposes.
Another model that shows promise in the prediction of noncake filtration
pressure drop is that of Leith and EllenbeckerflO] as modified by Koehler
and Leith.[11] In this model, the tube-sheet pressure drop is a function
of the clean fabric drag, the system hardware, and the cleaning energy.
Specifically:
\ [Pt + KlV'~ J(p<-Kiv/)2-4W0K,/K3] + KVV* (5.10)
where
P« = maximum static pressure achieved in the bag during
cleaning
KI = clean fabric resistance
V/ = face velocity
K2 = dust deposit flow resistance
Ka = bag cleaning efficiency coefficient
K,, = loss coefficient for the venturi at the inlet to the bag
Comparisons of laboratory data with pressure drops computed from Equa-
tion 5.10 [10,11] are in close agreement for a variety of dust/fabric com-
binations. The disadvantage of Equation 5.10 is that the constants K1?
{(2, and KB must be determined from laboratory measurements. The most
difficult one to determine is the KS value, which can only be found by
making measurements in a pilot-scale pulse-jet baghouse. A limitation of
laboratory measurements is that actual filtration conditions cannot always
be adequately simulated. For example, a redispersed dust may not have the
5-15
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same size distribution as the original dust, thereby yielding different values
of KI, K2, and K3 than would be measured in an operating baghouse.
5.2 Design Procedures
5.2.1 Gas-to-Cloth Ratio
The gas-to-cloth ratio is difficult to estimate from first principles. How-
ever, shortcut methods of varying complexity allow rapid estimation. De-
scriptions of three methods of increasing difficulty follow. For shaker and
reverse-air baghouses, the third method is best performed with publicly
. available computer model programs.
The methods outlined below pertain to conventional baghouses. Use of
electrostatic stimulation (an emerging technology) allows a higher gas-to-
cloth ratio at a given pressure drop; thus a smaller baghouse structure and
fewer bags are needed. This reduces total annual cost by up to 30%. Viner
and Locke [13] discuss cost and performance models for electrostatically
stimulated fabric filters.
5.2.1.1 Gas-to-Cloth Ratio From Similar Applications
After a fabric has been selected, the gas-to-cloth ratio can be determined
using Table 5.1. Column 1 shows the type of dust; column 2 shows the
gas-to-cloth ratios for woven fabric; and column 3 shows gas-to-cloth ratios
for felted fabrics. Notice that these values are all net gas-to-cloth ratios.
The net gas-to-cloth ratio is equal to the total actual volumetric flow rate
in cubic feet per minute divided by the net cloth area in square feet. This
ratio, in units of feet per minute, affects pressure drop and bag life. The net
cloth area is determined by dividing the gas-to-cloth ratio into the actual
cubic feet per minute flow of the exhaust gas stream. For an intermittent-
type baghouse that is shut down for cleaning, this is the total, or gross,
cloth area. However, for continuously operated filters, the area must be
increased to allow the shutting down of one or more compartments for
cleaning. Table 5.2 provides a guide for adjusting the net area to the gross
area, which determines the size of a continuously cleaned filter.
5-16
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Table 5.1: Gas-to-Cloth Ratios*'0
(actual ft3/min)/(ft2 of net cloth area)
Dust
Alumina
Asbestos
Bauxite
Carbon Black
Coal
Cocoa, Chocolate
Clay
Cement
Cosmetics
Enamel Frit
Feeds, Grain
Feldspar
Fertilizer
Flour
Fly Ash
Graphite
Gypsum
Iron Ore
Iron Oxide
Iron Sulfate
Lead Oxide
Leather Dust
Lime
Limestone
Mica
Paint Pigments
Paper
Plastics
Quartz
Rock Dust
Sand
Sawdust (Wood)
Silica
Slate
Soap, Detergents
Spices
Starch
Sugar
Talc
Tobacco
Zinc Oxide
•Reference [12]
Shaker/Woven
Reverse- Air/Woven
2.5
3.0
2.5
1.5
2.5
2.8
2.5
2.0
1.5
2.5
3.5
2.2
3.0
3.0
2.5
2.0
2.0
3.0
2.5
2.0
2.0
3.5
2.5
2.7
2.7
2.5
3.5
2.5
2.8
3.0
2.5
3.5
2.5
3.5
2.0
2.7
3.0
2.0
2.5
3.5
2.0
"Generally safe design values; application
eration of particle
size and grain loading.
Pulse Jet/Felt
Reverse-Air Felt
8
10
8
5
8
12
9
8
10
9
14
9
8
12
5
5
10
11
7
6
6
12
10
8
9
7
10
7
9
9
10
12
7
12
5
10
8
7
10
13
5
requires consid-
5-17
-------
Table 5.2: Approximate Guide to Estimate Gross Cloth Area
Net Cloth Area
(ft2)
1-4,000
4,001-12,000
12,001-24,000
24,001-36,000
36,001-48,000
48,001-60,000
60,001-72,000
72,001-84,000
84,001-96,000
96,001-108,000
108,001-132,000
132,001-180,000
above 180,001
Gross Cloth
(ft2)
Multiply by
n
n
n
n
n
n
n
n
n
n
n
n
Area
2
1.5
1.25
1.17
1.125
1.11
1.10
1.09
1.08
1.07
1.06
1.05
1.04
'Reference [14J
5-18
-------
5.2.1.2 Gas-to-Cloth Ratio From Manufacturer's Methods
Manufacturers have developed nomographs and charts that allow rapid
estimation of the gas-to-cloth ratio. Two examples are given below, one for
shaker-cleaned baghouses and the other for pulse-jet cleaned baghouses.
For shaker baghouses, Table 5.3 gives a factor method for estimating the
ratio. Ratios for several materials in different operations are presented, but
are modified by factors for particle size and dust load. Directions and an
example are included. Gas-to-cloth ratios for reverse-air baghouses would
be about the same or a little more conservative compared to the Table 5.3
values.
For pulse-jet baghouses, which normally operate at two or more times
the gas-to-cloth ratio of reverse-air baghouses, another factor rnethod[15]
has been modified with equations to represent temperature, particle size,
and dust load:
V = 2.878 A B T-°-23351-0-06021 (0.7471 + 0.0853 In D) (5.11)
where
V = gas-to-cloth ratio (ft/min)
A = material factor, from Table 5.4
B = application factor, from Table 5.4
T = temperature, (°F, between 50 and 275)
L = inlet dust loading (gr/ft3, between 0.05 and 100)
D = mass mean diameter of particle (/xm, between 3 and 100)
For temperatures below 50°F, use T = 50 but expect decreased accu-
racy; for temperatures above 275°F, use T = 275. For particle mass mean
diameters less than 3 /xm, the value of D is 0.8, and for diameters greater
than 100 /xm, D is 1.2. For dust loading less than 0.05 gr/ft3, use L = 0.05;
for dust loading above 100 gr/ft3, use L = 100.
5.2.1.3 Gas-to-Cloth Ratio From Theoretical/Empirical Equa-
tions
Shaker and reverse-air baghouses The system described by Equations
5.1 through 5.6 is complicated; however, numerical methods can be used to
5-19
-------
Table 5.3: Manufacturer's Factor Method for Estimating Gas-to-Cloth
Ratios for Shaker Baghouses
UAHHIAL
Caidboad
Feeds
Flow
Giam
Lealhei Dust
Tobacco
Supply AII
Kood, Oust. Chips
*ATIO
"oVtRATicm
I
2.3.4.5.6.7
2.3.4.5.6.7
2.3.4.5.6,7
1.7.8
1.4.6.7
13
1.6.7
3/1 tATIO
MATERIAL
Asbestos
Aluminum Oust
Fibiows Ual'l.
Cellulose Mat'l.
Gypsum
linefHyduled)
Peilile
RubbeiChem.
Sail
Saul*
lion Scale
Soda Ask
Talc
Machiiiin|0peialioij|,8
OPERATION
1.78
1.7.8
1.4.7.8
1.4.7.8
1,3.5.67
J.4.6.7
J.4.5.6
4.5.6.7.J
2.3.4.5,6.7
5.6.7.9.15
1.7.8
4.6.7
3.4.5.6.7
7 VI lATIO
Alumina
Obon Black
Cement
Coke
Ceiamic Piem.
Clay&BnckOusI
Coal
Kaolin
Limestone
Rock, Die Oust
Silica
Suta
OPERATION
2.3,4.5.6
4.5.6.7
3.4.5.6,7
2.3.5.6
4.5.6.7
2.4.6.12
2.3.6.7.12
4.5.7
2.3.4.5.8.7
2.3.4.5.6.7
2.3.4.5.6.7
3.4.5.6.7
~2~7l »Afib ~
MATERIAL
Ammonium Phos-
phate Peil.
Dialomaceous
Eailh
Dry Petiochem.
Dyes
Fly Ash
Uelal Powdeis
Plastics
Reims
Silicates
Slareh
Soaps
2.3.4.5. S. 7
4.5.6.7
2.3.4.5.6.7.14
2.3.4.5.6.7
10
2.3.4.5.6.7,14
2.3.4.5.6.7.14
2.3.4.5.6.7.14
2.3.4.5.6,7.14
U
3.4,5.6.7
MATERIAL
Activated Charcoal
Caibpn Black
Deteteenls
Metal Fumes,
(hides and
other Solid
Dispersed
Products
OPERATION
2,4.5.6.7
II. 14
2.4.5.6.7
10,11
CUTTING
CRUSHING • -
PUIVERI2ING-
MIXING - 4
SCREENING - 5
STORAGE '- 6
CONVEYING - 7
GRINDING -1
SHAKCOUT -9
FURNACE, fUMC - K)
REACTION FUME- II
DUMPING . I?
INTAKE CLEANING- 13
PROCESS - 14
•LASTING - 15
FINENESS FACTOR
MICRON SIZE FACTOR
_
50-100
10-50'
3-10
1-3
<1
I.I
7
C DUST LOAD FACTOR
loading GR. CU. FT. Facloc
Thii infermalien can>lilul« a guide for commonly •ncounlwcd lilualioni and thould not b* con-
>id«i*d a "hafd-and-fa>t" nil*. Air-to-cloln ratios arc d*p«nd*nl on dull loading, til* diilribulion,
parlicl* >hap« and "coh««iy«n««»" o< lh« dteotiUd dusl. Th«« condiliont mutt b« *valuol*d lor
tach opplkolion. The longer lh« interval b«lwt*n bag clraning >h« lower lh« air-to-clolh ratio
mini bt. Fimly-divided, uniformly tiled parlklet generally form more dente filler caket and re-
quire lower oir-to-clolh ratios than when larger particles ore interspersed with the fines. Stkky.
oily particles, regardless o« shape or siie, form dense filler cakes and require lower oir-lo-clolh
ratios.
1.2
9- 17
1.0
~ .95
18-40
.90
40
85
EXAMPIE: Foundry thakeoul unit handling 26000 CFM and collecting 3500 #/ hr. of land. The
particle diilribulion shows 90% greater1 than 10 microns. The air is to ohautt lo room
in winler, to atmosphere in summer.
3500 »Xi, H- 003$- 26000%£. X 7000°'/. = 15.7 ^
•Chan A — 3/1 ratio. Chart a = Factor 1.0. Chart C — 95; 3 « I « .95 - 2.9 oir
lo cloth ratio. 26000 -r- 2.9 = 9.000 sq. ft.
Reprinted with permission from Buffalo Forge Company Bulletin AHD-29.
5-20
-------
Table 5.4: Factors for Pulse-Jet Gas-to-Cloth Ratios *
A. Material
15°
Cake mix
Cardboard
dust
Cocoa
Feeds
Flour
Grain
Leather
dust
Sawdust
Tobacco
Factor
12
Asbestos
Buffing dust
Fibrous and
cellulosic
material
Foundry
shakeout
Gypsum
Lime
(hydrated)
Perlite
Rubber
chemicals
Salt
Sand
Sandblast
dust
Soda ash
Talc
10
Alumina
Aspirin
Carbon black
(finished)
Cement
Ceramic pigments
Clay and brick
dusts
Coal
Fluorspar
Gum, natural
Kaolin
Limestone
Perchlorates
Rock dust, ores
and minerals
Silica
Sorbic acid
Sugar
9.0
Ammonium
phosphate-
fertilizer
Cake
Diatomaceous
earth
Dry petro-
chemicals
Dyes
Fly ash
Metal powder
Metal oxides
Pigments
metallic and
synthetic
Plastics
Resins
Silicates
Starch
Stearates
Tannic acid
6.0*
Activated carbon
Carbon black
(molecular)
Detergents
Fumes and
other dispersed
products direct
from reactions
Powdeired milk
Soaps
B. Application Factor
Nuisance Venting
Relief of transfer points,
conveyors, packing stations, etc.
Product Collection
Air conveying-venting, mills,
flash driers, classifiers, etc.
Process Gas Filtration
Spray driers, kilns, reactors, etc.
1.0
0.9
0.8
•Reference [15]
"In general, physically and chemically stable materials.
6Also includes those solids that are unstable in their physical or chemical state due
to hygroscopic nature, sublimation, and/or polymerization.
5-21
-------
obtain an accurate solution. A critical weakness in baghouse modeling that
has yet to be overcome is the lack of a fundamental description of the bag
cleaning process. That is, to solve Equations 5.1 through 5.6, the value of
Wr (the dust load after cleaning) must be known. Clearly, there must be a
relationship between the amount and type of cleaning energy and the degree
of dust removal from a bag. Dennis et a/.[16] have developed correlations
for the removal of coal fly ash from woven fiberglass bags by shaker cleaning
and by reverse air cleaning. These correlations have been incorporated into
a computer program that generates the solution to the above system of
equations.[9,17,18] If one were to apply the correlations developed with coal
ash and woven glass fabrics to other dust/fabric combinations, the accuracy
of the results would depend on how closely that dust/fabric combination
mimicked the coal ash/woven glass fabric system.
Physical factors that affect the correlation include the particle size dis-
tribution, adhesion and electrostatic properties of the dust and fabric, and
fabric weave, as well as cleaning energy. More research is needed in this
area of fabric filtration.
The rigorous design of a baghouse thus involves several steps. First, the
design goal for average pressure drop (and maximum pressure drop, if neces-
sary) must be specified along with total gas flow rate and other parameters,
such as Se and Kj (obtained either from field or laboratory measurements).
Second, a face velocity is assumed and the number of compartments in
the baghouse is computed based on the total gas flow, face velocity, bag
size, and number of bags per compartment. (Typical compartments in the
U.S. electric utility industry use bags 1 ft in diameter by 30 ft long with
400 bags per compartment.) Standard practice is to design a baghouse
to meet the specified pressure drop when one compartment is off-line for
maintenance. The third step is to specify the operating characteristics of
the baghouse (i.e., filtration period, cleaning period, and cleaning mech-
anism). Fourth, the designer must specify the cleaning efficiency so that
the residual dust load can be estimated. Finally, the specified baghouse
design is used to establish the details for Equations 5.1 through 5.6, which
•are then solved numerically to establish the pressure drop as a function
of time. The average pressure drop is then computed by integrating the
instantaneous pressure drop over the filtration cycle and dividing by the
cycle time. If the computed average is higher than the design specifica-
tion, then the face velocity must be reduced and the procedure repeated.
If the computed average pressure drop is significantly lower than the de-
5-22
-------
sign specification, then the proposed baghouse was oversized and should be
made smaller by increasing the face velocity and repeating the procedure.
When the computed average pressure drop comes sufficiently close to the
assumed specified value, then the design has been determined. A complete
description of the modeling process can be found in the reports by Dennis
et a/.[16,18] A critique on the accuracy of the model is presented by Viner
et a
Pulse-jet baghouses The overall process of designing a pulse jet bag-
house is actually simpler than that required for a reverse-air or shaker
baghouse if the baghouse remains on-line for cleaning. The first step is
to specify what the desired average tube-sheet pressure drop should be.
Second, the operating characteristics of the baghouse must be established
(e.g., on-line time, cleaning energy). Third, the designer must obtain values
for the coefficients in either Equation 5.9 or Equation 5.10 from field, pilot
plant, or laboratory measurements. Fourth, a value is estimated for the
face velocity and the appropriate equation (Equation 5.8 or 5.10) is solved
for the pressure drop as a function of time for the duration of the filtration
cycle. This information is used to calculate the cycle average pressure drop.
If the calculated pressure drop matches the specified pressure drop, then
the procedure is finished. If not, then the designer must adjust the face
velocity and repeat the procedure.
5.2.2 Pressure Drop
Pressure drop for the bags can be calculated rigorously from the equa-
tions given in the preceding section if values for the various parameters are
known. Frequently they are not known. For quick estimation, a maximum
pressure drop of 5 to 10 in. H2O across the baghouse and 10 to 20 in. H2O
across the entire system can be assumed if it contains much ductwork.
A comparable form of Equations 5.1 and 5.3 that may be used for pres-
sure drop across the fabric in a shaker or reverse-air baghouse is:
AP = SeV + K,dV2e (5.12)
5-23
-------
where
AP = pressure drop (in. H2O)
Se = effective residual drag of the fabric [in. H2O/(ft/min)]
V = superficial face velocity or gas-to-cloth ratio (ft/min)
K2 = specific resistance coefficient of the dust
{[in. H20/(ft/min)]/(lb/ft2)}
Cj = inlet dust concentration (lb/ft3)
9 = filtration time (min)
Although there is much variability, values for Se may range from about
0.2 to 2 in. H20/(ft/min) and for K2 from 1-2 to 30-40 [in. H20/(ft/min)]/
(lb/ft2). Typical values for coal fly ash are about 1 to 4. Inlet concentrations
vary from less than 0.05 gr/ft3 to more than 100 gr/ft3, but a more nearly
typical range is.from about 0.5 to 10 gr/ft3. Filtration times may range
from about 20 to 90 minutes for continuous duty baghouses, but 30 to 60
minutes is more frequently found. For pulse-jet baghouses, use Equations
5.8 and 5.9 to estimate AP, after substituting dVd for W0 and (PE)Alo for
S«V.
5.2.3 Particle Characteristics
Particle size distribution and adhesiveness are the most important particle
properties that affect design procedures. Smaller particle sizes can form
a denser cake, which increases pressure drop. As shown in Table 5.3 and
Equation 5.11, the effect of decreasing average particle size is a lower ap-
plicable gas-to-cloth ratio.
Adhering particles, such as oily residues or electrostatically active plas-
tics, may require installing equipment that injects a precoating material
onto the bag surface, which acts as a buffer that traps the particles and
prevents them from blinding or permanently plugging the fabric pores. In-
formed fabric selection may eliminate electrostatic problems.
5.2.4 Gas Stream Characteristics
Moisture and corrosives content are the major gas stream characteristics
requiring design consideration. The baghouse and associated ductwork
5-24
-------
should be insulated and possibly heated if condensation may occur. Both
the structural and fabric components must be considered, as either may be
damaged. Where structural corrosion is likely, stainless steel substitution
for mild steel may be required, provided that chlorides are not present.
(Most austenitic stainless steels are susceptible to chloride corrosion.)
5.2.4.1 Temperature
The temperature of the pollutant stream to be cleaned must be above and
remain above the dew point of any "condensables in the stream. If the
temperature is high and it can be lowered without approaching the dew
point, spray coolers or dilution air can be used to drop the temperature so
that temperature limits of the fabric will not be exceeded. The additional
cost of a precooler will have to be weighed against the higher cost of bags
with greater temperature resistance. The use of dilution air to cool the
stream also constitutes a tradeoff between a less expensive fabric and a
larger filter necessary to accommodate the additional volume of the dilution
air. Generally, precooling would not be necessary if fabric that will handle
the temperature and the chemical action of the pollutant stream is available.
(Costs for spray chambers, quenchers, and other precoolers are found in
the "Precoolers" section of the Manual.) Table 5.5 lists several of the
fabrics in current use and provides information on temperature limits and
chemical resistance. The column labeled "Flex Abrasion" indicates the
fabric's suitability for cleaning by mechanical shakers.
5.2.4.2 Pressure
Standard fabric filters can be used in pressure or vacuum service but only
within the range of about ±25 inches of water. Because of the sheet metal
construction of the house, they are not generally suited for more severe
service. However, for special applications, high-pressure shells can be built.
5-25
-------
Table 5.5: Properties of Leading Fabric Materials*
Fabric
Cotton
Creslan*
Dacron6
Temp,
opa
180
250
275
Acid
Resistance
Poor
Good in mineral
acids
Good in most
Alkali
Resistance
Very good
Good in weak alkali
Good in weak alkali;
Flex
Abrasion
Very good
Good to very
good
Very good
Dynele
160
mineral acids; dis-
solves partially in
concentrated
HaSO4
Little effect even
at high concentra-
tion
fair in strong alkali
Little effect even in
high concentration
Fair to good
Fiberglas''
Filtron8
Gore-Tex'
Nomex*
Nylon"
Orion'
Polypropylene
Teflon4
Wool
500
270
Depends
on
backing
375
200
260
200
450
200
Fair to good
Good to excellent
Depends on back-
ing
Fair
Fair
Good to excellent
in mineral acids
Excellent
Inert except to flu-
orine
Very good
Fair to good
Good
Depends on backing
Excellent at low
temperature
Excellent
Fair to good in weak
alkali
Excellent
Inert except to triflu-
oride, chlorine, and
molten alkaline met-
als
Poor
Fair
Good to very
good
Fair
Excellent
Excellent
Good
Excellent
Fair
Fair to good
•Reference [20]
"Maximum continuous operating temperatures recommended by the Industrial Gas
Cleaning Institute.
6 American Cyanamid registered trademark.
eDu Pont registered trademark.
''Owens-Corning Fiberglas registered trademark.
'W. W. Criswell Div. of Wheelabrator-Fry, Inc., trade name.
. L. Gore and Co., registered trademark.
5-26
-------
5.2.5 Equipment Design Considerations
5.2.5.1 Pressure or Suction Housings
The location of the baghouse with respect to the fan in the gas stream af-
fects the capital cost. A suction-type baghouse, with the fan located on the
downstream side of the unit, must withstand high negative pressures and
therefore must be more heavily constructed and reinforced than a baghouse
located downstream of the fan (pressure baghouse). The negative pressure
in the suction baghouse can result in outside air infiltration, which can re-
sult in condensation, corrosion, or even explosions if combustible gases are
being handled. In the case of toxic gases, this inward leakage can have an
advantage over the pressure-type baghouse, where leakage is outward. The
main advantage of the suction baghouse is that the fan handling the pro-
cess stream is located at the clean-gas side of the baghouse. This reduces
the wear and abrasion on the fan and permits the use of more efficient fans
(backward-curved blade design). However, because for some designs the
exhaust gases from each compartment are combined in the outlet manifold
to the fan, locating compartments with leaking bags may be difficult and
adds to maintenance costs. Pressure-type baghouses are generally less ex-
pensive because the housing must only withstand the differential pressure
across the fabric. In some designs the baghouse has no external housing.
Maintenance also is reduced because the compartments can be entered and
leaking bags can be observed while the compartment is in service. With
a pressure baghouse, the housing acts as the stack to contain the fumes
with the subsequent discharge at the roof of the structure, which makes it
easier to locate leaking bags. The main disadvantage of the pressure-type
baghouse is that the fan is exposed to the dirty gases where abrasion and
wear on the fan blades may become a problem.
5.2.5.2 Standard or Custom Construction
The design and construction of baghouses are separated into two groups,
standard and custom,[14] which are further separated into low, medium,
and high capacity. Standard baghouses are predesigned and factory built
as complete off-the-shelf units that are shop-assembled and bagged for
low-capacity units (hundreds to thousands of acfm throughput). Medium-
capacity units (thousands to less than 100,000 acfm) have standard designs,
5-27
-------
are shop-assembled, may or may not be bagged, and have separate bag
compartment and hopper sections. High-capacity baghouses (larger than
50,000 or 100,000 acfm) can be designed as shippable modules requiring
only moderate field assembly. These modules may have bags installed and
can be shipped by truck or rail. Upon arrival, they can be operated singly
or combined to form units for larger-capacity applications. Because they
are preassembled, field labor for installation is less costly.
The custom baghouse, also high capacity, is designed for a specific appli-
cation and is usually built to the specifications prescribed by the customer.
Generally, these units are much larger than standard baghouses. For ex-
ample, many are used on power plants. The cost of the custom baghouse
is much higher per square foot of fabric because it is not an off-the-shelf
item and requires special setups for manufacture and expensive field labor
for assembly upon arrival. The advantages of the custom baghouse are
many and are usually directed towards ease of maintenance, accessibility,
and other customer preferences. In some standard baghouses, a complete
set of bags must be replaced in a compartment at one time because of the
difficulty in locating and replacing single leaking bags, whereas in custom
baghouses, single bags are accessible and can be replaced one at a time as
leaks develop.
5.2.5.3 Filter Media
The type of filter material used in baghouses is dependent on the specific
application in terms of chemical composition of the gas, operating temper-
ature, dust loading, and the physical and chemical characteristics of the
particulate. A variety of fabrics, either felted or woven, is available and the
selection of a specific material, weave, finish, or weight is based primarily
on past experience. The type of yarn (filament, spun, or staple), the yarn
diameter, and twist are also factors in the selection of suitable fabrics for
a specific application. For some difficult applications, Gore-Tex, a polyte-
trafluoroethylene (PTFE) membrane laminated to a fabric (felt or woven)
may be used. Because of the violent agitation of mechanical shakers, spun
or heavy weight staple yarn fabrics are commonly used with this type of
cleaning, while lighter weight filament yarn fabrics are used with reverse-air
cleaning.
The type of material will limit the maximum operating gas temperature
5-28
-------
for the baghouse. Cotton fabric has the least resistance to high temper-
atures (about 180°F), while fiberglass has the most (about 500°F). The
temperature of the exhaust-gas stream must be well above the dew point
of any of its contained condensables, as liquid particles will usually plug the
fabric pores quickly. However, the temperature must be below the maxi-
mum limit of the fabric in the bags. These maximum limits are given in
Table 5.5.
5.3 Estimating Total Capital Investment
Total capital investment includes costs for the baghouse structure, the ini-
tial complement of bags, auxiliary equipment, and the usual direct and
indirect costs associated with installing or erecting new structures. These
costs are described below. (Costs for improving baghouse performance with
electrical enhancement are not discussed in this section, but are mentioned
in the example problem.)
5.3.1 Equipment Cost
5.3.1.1 Bare Baghouse Costs
Cost correlations for six types of baghouses are presented. These six types,
five of which are preassembled and one, field-assembled, are outlined in
Table 5.6.
Each figure gives costs for the filter without bags and additional costs
for stainless steel construction and for insulation. All curves are based on
a number of actual quotes. A least squares line has been fitted to the
quotes and the line's equation is given. In most cases these lines should
not be extrapolated in either direction. The reader should not be surprised
if he obtains quotes that differ from these curves by as much as ±25%.
Significant savings can be obtained by soliciting multiple quotes. All units
include inlet and exhaust manifolds, supports, platforms, handrails, and
hopper discharge devices. The indicated prices are flange to flange. The
scales on both axes change from one figure to another to accommodate the
5-29
-------
Table 5.6: Scope of Cost Correlations
Baghouse Type Figure No.
Intermittent
Continuous
Continuous
Continuous
Continuous
Preassembled units
Shaker
Shaker
Pulse-jet (common housing)
Pulse-jet (modular)
Reverse- air
5.2
5.3
5.4
5.5
5.6
Field-assembled units
Continuous Any method 5.7
differing gas flow ranges over which the various types of baghouses operate.
The 304 stainless steel add-on cost is used when such construction is
necessary to prevent the exhaust gas stream from corroding the interior of
the baghouse. Stainless steel is substituted for all metal surfaces that are
in contact with the exhaust gas stream.
Insulation costs are for 3 inches of shop-installed glass fiber encased in a
metal skin. One exception is the custom baghouse, which has field-installed
insulation. Costs for insulation include only the flange-to-flange baghouse
structure on the outside of all areas in contact with the exhaust gas stream.
Insulation for ductwork, fan casings, and stacks must be calculated sepa-
rately as discussed later.
The first baghouse type is the intermittent service baghouse cleaned by
a mechanical shaker. This baghouse is shut down and cleaned at conve-
nient times, such as the end of the shift or end of the day. Although few
units are sold, they are applicable for operations that require infrequent
cleaning. Figure 5.2 presents the unit cost with price in dollars plotted
against the gross square feet of cloth required. [21] Because intermittent
service baghouses do not require an extra compartment for cleaning, gross
and net fabric areas are the same. The plot is linear because baghouses
are made up of modular compartments and thus have little economy of
5-30
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til
Caution: Do not extrapolate.
Stainless steel add on
4 6 8 10 12 14 16
Gross Cloth Area (1000ft2)
18
Figure 5.2: Equipment costs for intermittent shaker filters
scale. Because of the modular construction, the price line should not be
extrapolated downward.
Figure 5.3 presents similar costs for a continuously operated baghouse
cleaned by mechanical shaker.[21,22] Again, price is plotted against the
gross cloth area in square feet. As in Figure 5.2, the units are modular
in construction. Costs for these units, on a square foot basis, are higher
because of increased complexity and generally heavier construction.
The third and fourth types are common-housing pulse jets and modular
pulse jets. The latter are constructed of separate modules that may be ar-
ranged for off-line cleaning, and the former have all bags within one housing.
The costs for these units are shown in Figures 5.4 and 5.5, respectively.[2i]
Note that in the single-unit (common-housing) pulse jet, for the range
shown, the height and width of the unit are constant and the length in-
creases; thus, for a different reason than that for the modular units dis-
cussed above, the cost increases linearly with size. Because the common
5-31
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Caution: Do not extrapolate.
Stainless steel add on
10 20 30 40 50 60 70 30 9(
Gross Goth Area (1000ft2)
Source: ETS, Inc.; Fuller Co.
Figure 5.3: Equipment costs for continuous shaker filters.
5-32
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Caution: Do not extrapolate.
Cost without bags: Z
Stainless steel add on -
Insulation add onH
6 8 10 12 14 16
Gross Cloth Area (1000ft2)
Source: ETS, Inc.
Figure 5.4: Equipment costs for pulse-jet filters (common housing).
5-33
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Caution: Do not extrapoic;«.
Stainleu ateal add on
Inaulatlon add on- -
24 6 8 10 12 14 16 1
Gross Cloth Area (1000ft^
Source: ETS. Inc.
Figure 5.5: Equipment costs for pulse-jet filters (modular).
5-34
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Caution: Do not extrapolate.
I 102030405060708090
Gross doth Area (1000 ft ^
Source: ETS. Inc.
Figure 5.6: Equipment costs for reverse-air filters.
housing is relatively inexpensive, the stainless steel add-on is proportion-
ately higher than for modular units. Added material costs and setup and
labor charges associated with the less workable stainless steel account for
most of the added expense.
Figure 5.6 shows the costs for the reverse-air baghouses.[21] Again, the
construction is modular. The final type is the custom baghouse which,
because of its large size, must be field assembled. It is often used on power
plants, steel mills, or other applications too large for the factory-assembled
baghouses. Prices for custom units are shown in Figure 5.7. [21]
5.3.1.2 Bag Costs
Table 5.7 gives the price per square foot of bags by type of fabric and by
type of cleaning system used. Actual quoted prices may vary by ±10%
from the values in the table. In calculating the cost, the gross area as
5-35
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Caution: Do not extrapolate.
100 200 300
Gross doth Area (1000ft2)
Source: ETS, Inc.
Figure 5.7: Equipment costs for custom-built filters.
5-36
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Table 5.7: Bag Prices
(3rd quarter 1986 $/ft2)
Type
Type of Cleaning
Pulse jet, TR*
Pulse jet, BBR
Shaker, Strap top
Shaker, Loop top
Reverse air with rings
Bag Diameter
(inches)
4-1/2
6
4-1/2
6
to
to
to
to
5
5
8
5-1/8
8
5-1/8
8
11-1/2
Reverse air w/o rings
8
11-1/2
PE
0.59
0.43
0.37
0.32
0.45
0.43
0.46
0.47
0.32
0.32
PP
0.61
0.44
0.40
0.33
0.48
0.45
NA
NA
NA
NA
NO
1
1
1
1
1
1
I
1
1
1
.88
.56
.37
.18
.28
.17
.72
.69
.20
.16
of Material"
HA
0.92
0.71
0.66
0.58
0.75
0.66
NA
NA
NA
NA
FG
1.29
1.08
1.24
0.95
NA
NA
0.99
0.76
0.69
0.53
CO
NA
NA
NA
NA
0.44
0.39
NA
NA
NA
NA
TF
9.05
6.80
8.78
6.71
NA
NA
NA
NA
NA
NA
NA = Not applicable.
"Materials:
PE = 16-08 polyester FG = 16-oi fiberglass with 10% Teflon
PP = 16-o* polypropylene CO — 9-oi cotton
NO = 14-01 nomex TF = 22-o» Teflon felt
HA = 16-oz homopolymei acrylic
*Bag removal methods:
TR = Top bag removal (snap in)
BBR = Bottom bag removal
NOTE: For pulse-jet baghouses, all bags are felts except for the fiberglass, which is woven.
For bottom access pulse jets, the mild steel cage price for one cage can be calculated from
the single-bag fabric area using:
$ = 4.941 + 0.163 ft1 in 50 cage lots
$ = 4.441 + 0.163 ft2 in 100 cage lots
8 = 3.941 + 0.163 ft2 in 500 cage lots
These costs apply to 4-1/2 inch or 5-5/8 inch diameter, 8 foot and 10 foot cages made of
11 gauge mild steel and having 10 vertical wires and "Roll Band" tops. For flanged tops,
add $1 per cage. If flow control Venturis are used (as they are in about half of the pulse-jet
manufacturers' designs), add $5 per cage. For stainless steel cages use:
$ = 23.335 + 0.280 ft2 in 50 cage lots
$ = 21.791 + 0.263 ft2 in 100 cages lots
S = 20.564 + 0.248 ft2 in 500 cage lots
For shakers and reverse air baghouses, all bags are woven. All prices are for finished bags,
and prices can vary from one supplier to another. For Gore-Tex bag prices, multiply base
fabric price by factors of 3 to 4.5.
Source: ETS, Inc.[21] _ __
-------
determined from Table 5.2 should be used. Gore-Tex fabric costs are a
combination of the base fabric cost and a premium for the PTFE laminate
and its application. As fiber market conditions change, the costs of fabrics
relative to each other also change. The bag prices are based on typical
fabric weights, in ounces/square yard, for the fabric being priced. Sewn-in
snap rings are included in the price, but other mounting hardware, such as
clamps or cages, is an added cost.
5.3.1.3 Auxiliary Equipment
The auxiliary equipment depicted in Figure 5.1 is discussed elsewhere in
the Manual. Because hoods, ductwork, precoolers, cyclones, fans, motors,
dust removal equipment and stacks are common to many pollution con-
trol systems, they are (or will be) given extended treatment in separate
chapters.
5.3.2 Total Purchased Cost
The total purchased cost of the fabric filter system is the sum of the costs of
the baghouse, bags, auxiliary equipment, instruments and controls; and of
taxes and freight. The last three items generally are taken as percentages
of the estimated total cost of the first three items. Typical values, from
Chapter 2, are 10% for instruments and controls, 3% for taxes, and 5% for
freight.
Bag costs can vary from less than 15% to more than 100% of bare
baghouse cost, depending on type of fabric required. This situation makes
it inadvisable to estimate total purchased cost without considering both
costs, and prevents effective use of factors to estimate a single cost for the
baghouse and bags.
5.3.3 Total Capital Investment
Using Chapter 2 methodology, the total capital investment (TCI) for most
baghouses is estimated from a series of factors applied to the purchased
equipment cost to obtain direct and indirect installation costs. The TCI is
5-38
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the sum of these three costs (i.e., purchased equipment and direct and indi-
rect installation costs). The required factors are given in Table 5.8. Because
bag costs can have such a large effect on the total purchased equipment cost,
the factors may cause overestimation of total capital investment when ex-
pensive bags are used. Using stainless steel components may also cause
overestimation. Because baghouses may vary from small units installed
within existing buildings to large, separate structures, specific factors for
site preparation or for buildings are not given. However, costs for buildings
may be obtained from such references as Means Square Foot Costs 1986.[23]
Land, working capital, and off-site facilities are excluded from the table, as
they are not normally required. For very large installations, however, they
may be needed and would be estimated on an as-needed basis.
The factors given in Table 5.8 are for average installation conditions.
Considerable variation may be seen with other-than-average installation
circumstances. Moreover, the Table 5.8 factors may be too large for "pack-
aged" fabric filters—those pre-assembled baghouses that consist of the com-
partments, bags, waste gas fan and motor, and instruments and controls.
Because these packaged units require very little installation, their installa-
tion costs would be lower (20-25% of the purchased equipment cost).
5.4 Estimating Total Annual Costs
5.4.1 Direct Annual Cost
Direct annual costs include operating and supervisory labor, operating ma-
terials, replacement bags, maintenance (labor and materials), utilities, and
dust disposal. Most of these costs are discussed individually below. They
vary considerably with location and time, and, for this reason, should be
obtained to suit the specific baghouse system being costed. For example,
current labor rates may be found in such publications as the Monthly La-
bor Review, published by the U.S. Department of Labor, Bureau of Labor
Statistics.
5-39
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Table 5.8: Capital Cost Factors for Fabric Filters"
Cost Item Factor
Direct costs
Purchased equipment costs
Fabric filter (EC) + bags + auxiliary equipment As estimated, A
Instrumentation 0.10 A
Sales taxes 0.03 A
Freight o.05_A
Purchased Equipment Cost, PEC B = 1.18 A
Direct installation costs
Foundations & supports 0.04 B
Handling & erection 0.50 B
Electrical 0.08 B
Piping 0.01 B
Insulation for ductwork6 0.07 B
Painting0 Q.Q2 B
Direct installation cost Q.72 B
Site preparation As required, SP
Buildings As required, Bldg.
Total Direct Cost 1.72 B + SP + B7d~gT
Indirect Costs (installation)
Engineering 0.10 B
Construction and field expense 0.20 B
Contractor fees 0.10 B
Start-up 0.01 B
Performance test 0.01 B
Contingencies 0.03 B
Total Indirect Cost, 1C Q.45 B
Total Capital Investment = DC + 1C 2.17 B + SP~+ Bldg!
"Reference [24]
»If ductwork dimensions have been established, cost may be estimated based on $10 to $12/ft* (fourth
quarter 1988) of surface for field application. Fan housings and stacks may also be insulated.[2l]
'The increased use of special coatings may increase this factor to 0.08B or higher.[25]
5-40
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5.4.1.1 Operating and Supervisory Labor
Typical operating labor requirements are 2 to 4 hours per shift for a wide
range of filter sizes.[24] Small or well-performing units may require less
time, while very large or troublesome units may require more. Supervisory
labor is taken as 15% of operating labor.
5.4.1.2 Operating Materials
Operating materials are generally not required for baghouses. An exception
is the use of precoat materials injected on the inlet side of the baghouse
to provide a protective dust layer on the bags when sticky or corrosive
particles might harm them. Adsorbents may be similarly injected when
the baghouse is used for simultaneous particle and gas removal,, Costs for
these materials should be included on a dollars-per-mass basis (e.g., dollars
per ton).
5.4.1.3 Maintenance
Maintenance labor varies from 1 to 2 hours per shift.[24] As with operating
labor, these values may be reduced or exceeded depending on the size and
operating difficulty of a particular unit. Maintenance materials costs are
assumed to be equal to maintenance labor costs.[24]
5.4.1.4 Replacement Parts
The major replacement part items are filter bags, which have a typical
operating life of about 2 years. The following formula is used for computing
the bag replacement cost:
CRCB = (CB + CL) x CRFB (5.13)
5-41
-------
where
= bag capital recovery cost ($/year)
= initial bag cost including taxes and freight ($)
= bag replacement labor ($)
= capital recovery factor whose value is a function of
the annual interest rate and the useful life of the bags
(For instance, for a 10% interest rate and a 2-year life,
CRFB = 0.5762.)
The bag replacement labor cost (C^) will depend on such factors as
the number, size, and type of the bags; their accessibility; how they are
connected to the baghouse tube-sheet; etc. For example, in a reverse-air
baghouse it would probably take from 10 to 20 man-minutes to change an
8 inch by 24 foot bag that is clamped in place. This bag has a filtering
surface area of approximately 50 ft2. If the replacement labor rate were
$21.12/h (including overhead), CL would be from $0.07 to $0.14/ft2 of bag
area. As Table 5.7 shows, for some bags (e.g., cotton), this range of C^
would constitute a significant fraction of the purchased cost. For pulse jets,
replacement time would be about 5 to 10 man-minutes for a 5 inch by 10
foot bag in a top-access baghouse. These bag replacement times are based
on changing a minimum of an entire module and on having typical baghouse
designs. Times would be significantly longer if only a few bags were being
replaced or if the design for bag attachment or access were atypical.
This method treats the bags as an investment that is amortized over the
useful life of the bags, while the rest of the control system is amortized over
its useful life (typically 20 years; see Subsection 5.4.2). Values of CRFB for
bag lives different from 2 years can be calculated from Equation 2.3.
5.4.1.5 Electricity
Power is required to operate system fans and cleaning equipment. Fan
power for primary gas movement can be calculated from Equation 2.7.
After substituting into this equation a combined fan-motor efficiency of
0.65 and a specific gravity of 1.000, we obtain:[26]
Power/on = 0.000181
-------
where
Power/on = fan power requirement (kWh/yr)
Q = system flow rate (acfm)
AP = system pressure drop (in. H2O)
9 = operating time (h/yr)
Cleaning energy for reverse-air systems can be calculated from the num-
ber of compartments to be cleaned at one time (usually one, sometimes
two), and the reverse gas-to-cloth ratio (from about one to two times, the
forward gas-to-cloth ratio). Reverse-air pressure drop varies up to 6 or 7
in. H20 depending on location of the fan pickup (before or after the main
system fan).[27] The reverse-air fan generally runs continuously.
Typical energy consumption in kWh/yr for a shaker system operated
8,760 h/yr can be calculated from:[5]
P = 0.0534 (5.15)
where
A = gross fabric area (ft2)
5.4.1.6 Fuel
If the baghouse or associated ductwork is heated to prevent condensation,
fuel costs should be calculated as required. These costs can be significant,
but may be difficult to predict. For methods of calculating heat transfer
requirements, see Perry.[26]
5.4.1.7 Water
Cooling process gases to acceptable temperatures for fabrics being used
can be done by dilution with air, evaporation with water, or heat exchange
with normal equipment. The last two cases require consumption of plant
water, although costs are not usually significant. Section 4.4 of Chapter 4
provides information on estimating cooling-water costs.
5-43
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5.4.1.8 Compressed Air
Pulse-jet filters use compressed air at pressures of about 60 to 100 psig.
Typical consumption is about 2 scfm/1,000 cfm of gas filtered.[5] For ex-
ample, a unit filtering 20,000 cfm of gas uses about 40 scfm of compressed
air for each minute the filter is operated.
5.4.1.9 Dust Disposal
If collected dust cannot be recycled or sold, it must be landfilled or disposed
of in some other manner. Disposal costs are site-specific, but they may
typically run $20 or $30 per ton exclusive of transportation (see Section
2.4, Chapter 2).
5.4.2 Indirect Annual Cost
These include such costs as capital recovery, property taxes, insurance,
administrative costs ("G&A"), and overhead. The capital recovery cost
is based on the equipment lifetime and the annual interest rate employed.
(See Chapter 2 for a thorough discussion of the capital recovery cost and
the variables that determine it.) For fabric filters, the system lifetime varies
from 5 to 40 years, with 20 years being typical.[24] However, this does not
apply to the bags, which usually have much shorter lives. (See Section
5.4.1.) Therefore, as Chapter 2 suggests, when figuring the system capital
recovery cost, one should base it on the installed capital cost less the cost
of replacing the bags (i.e., the purchased cost of the bags plus the cost of
labor necessary to replace them). In other words:
CRC, = [TCI -CB- C^CRF, (5.16)
where
CRC, = capital recovery cost for fabric filter system ($/yr)
TCI = total capital investment ($)
CB = initial cost of bags including taxes and freight ($)
G£ = labor cost for replacing bags ($)
CRF, = capital recovery factor for fabric filter system (defined
in Chapter 2).
5-44
-------
For example, for a 20-year system life and a 10% annual interest rate,
the CRF. would be 0.1175.
As Chapter 2 suggests, the suggested factor to use for property taxes, in-
surance, and administrative charges is 4% of the TCI. Finally, the overhead
is calculated as 60% of the sum of operating, supervisory, and maintenance
labor, and maintenance materials.
5.4.3 Recovery Credits
For processes that can reuse the dust collected in the baghouse or that can
sell the dust in a local market, such as fly ash sold as an extender for paving
mixes, a credit should be taken. As used below, this credit (RC) appears
as a negative cost.
5.4.4 Total Annual Cost
Total annual cost for owning and operating a fabric filter system is the sum
of the components listed in Sections 5.4.1 through 5.4.3, i.e.:
TAG = DC + IC-RC (5.17)
where
TAG = total annual cost ($)
DC = direct annual cost ($)
1C = indirect annual cost ($)
RC = recovery credits (annual) ($)
5.4.5 Example Problem
Assume a baghouse is required for controlling fly ash emissions from a coal-
fired boiler. The flue gas stream is 50,000 acfm at 325°F and has an ash
loading of 4 gr/ft3. Analysis of the ash shows a mass median diameter of
7 fj.m. Assume the baghouse operates for 8,640 h/yr (360 d).
The gas-to-cloth ratio (G/C) can be taken from Table 5.1 as 2.5, for
woven fabrics in shaker or reverse-air baghouses, or 5, for felts used in
5-45
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pulse-jet baghouses. If a factor method were used for estimating G/C,
Table 5.3 for shakers would yield the following values: A = 2, B = 0.9, and
C = 1.0. The gas-to-cloth ratio would be:
2 x 0.9 x 1.0 = 1.8.
This value could also be used for reverse-air cleaning. For a pulse-jet unit,
Table 5.4 gives a value of 9.0 for factor A and 0.8 for factor B. Equation
5.11 becomes:
V = 2.878 x 9.0 x 0.8(275)-°-233B(4)-°-06021(0.7471 + 0.0853 In 7) (5.18)
= 4.69
Because this value is so much greater than the shaker/reverse-air G/C,
we conclude that the pulse-jet baghouse would be the least costly design.2
Assume the use of on-line cleaning in a common housing structure and, due
to the high operating temperature, the use of glass filter bags (see Table
5.5). At a gas-to-cloth ratio of 4.69, the fabric required is3
50,000 acfm/4.69 fpm = 10,661 ft2.
is:
From Figure 5.4, the cost of the baghouse ("common housing" design)
Cost = 9,688 + 5.552(10,661) = $68,878 (5.19)
Insulation is required. The insulation add-on cost from Figure 5.4 is:
Cost = 1,428 + 0.931(10,661) = $11,353 (5.20)
From Table 5.7, bag costs are $1.24/ft2 for 5-1/8 inch diameter glass
fiber, bottom removal bags. Total bag cost is
10,661 ft2 x $1.24/ft2 = $13,220.
For 10 ft long cages,
'This conclusion is based on the inference that a much higher G/C would yield lower
capital and, in turn, annual costs. However, to make a more rigorous selection, we would
need to calculate and compare the total annual costs of all three baghouse designs (assum-
ing all three are technically acceptable). The reader is invited to make this comparison.
Further discussion of the effects of G/C increases, and accompanying pressure drop in-
creases, on overall annual costs will be found in Reference 27.
'This is the total (gross) bag area required. No bag area adjustment factor has been
applied here, because this is a common housing pulse jet unit that is cleaned continuously
during operation. Thus, no extra bag compartment is needed, and the gross and net bag
areas are equal.
5-46
-------
fabric area per cage = 5| in./12 in./ft XTT x 10 ft
= 13.42 ft2.
The number of cages = 10,661 fta/13.42 ft2
= 795 cages (rounded up to next integer).
From Table 5.7, individual cage cost is
3.941 + 0.163(13.42ft2) = $6.128.
Total cage cost is
795 cages x $6.128/cage = $4,872.
Assume the following auxiliary costs have been estimated from data in
other parts of the Manual:
Ductwork $14,000
Fan 14,000
Motor 7,000
Starter 3,500
Dampers 7,200
Compressor 6,000
Screw conveyor 4,000
Stack _ Tj?99_
Total $62,700"
Direct costs for the fabric filter system, based on the factors in Table 5.8,
are given in Table 5.9. (Again, we assume site preparation and buildings
costs to be negligible.) Total capital investment is $412,000. Table 5.10
gives the direct and indirect annual costs, as calculated from the factors
given in Section 5.4.1. For bag replacement labor, assume 10 min per bag
for each of the 795 bags. At a maintenance labor rate of $21.12 (including
overhead), the labor cost is $2,809 for 133 h. The bags are assumed to
be replaced every 2 yr. The replacement cost is calculated using Equation
5.13.
Pressure drop (for energy costs) can be calculated from Equations 5.8
and 5.9, with the following assumed values:
,g in. H,0/(ft/min)
-
5-47
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Table 5.9: Capital Costs for Fabric Filter System
Example Problem
Cost Item Cost
Direct Costs
Purchased equipment costs
Fabric filter (with insulation)(EC) $80,231
Bags and cages 18,092
Auxiliary equipment 62,700
Sum = A $161,623
Instrumentation, 0.1A 16,102
Sales taxes, 0.03A 4,831
Freight, 0.05A 8,051
Purchased equipment cost, B $190,007
Direct installation costs
Foundation and supports, 0.04B 7,600
Handling and erection, 0.50B 95,004
Electrical, 0.08B 15,201
Piping, 0.01B 1,900
Insulation for ductwork, 0.07B 13,300
Painting, 0.02B 3,800
Direct installation cost $136,805
Site preparation —
Facilities and buildings —
Total Direct Cost $326,812
Indirect Costs (installation)
Engineering, 0.10B 19,001
Construction and field expenses, 0.20B 38,001
Contractor fees, 0.10B 19,001
Start-up, 0.01B 1,900
Performance test, 0.01B 1,900
Contingencies, 0.03B 5,700
Total Indirect Cost $85,503
Total Capital Investment (rounded) $412,000
5-48
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Table 5.10: Annual Costs for Fabric Filter System
Example Problem
Cost Item
Calculations
Cost
Direct Annual Costs, DC
Operating labor
Operator
Supervisor
Operating materials
Maintenance
Labor
Material
Replacement parts, bags
Utilities
Electricity
Compressed air
(dried and filtered)
Waste disposal
Total DC
Indirect Annual Costs, 1C
Overhead
6 h v 360 days v $12
3ay x yr x 'h
15% of operator = .15 x 25,920
3h ., 360 days
day x yr x h
100% of maintenance labor
[2,809 + (13,220 x 1.08a)] x 0.5762
0.000181 x 50,000acfm x 10.3 in. H3O x
8.640 h $006
yr x IWTT
2 scfrn
8,640 h
at $20/ton on-site for essentially 100% collection
~A3~ x T fififi^ x —*=i=— x —^S x
2,000 Ib
8.640 h 1 ton
x X
Administrative charges
Property tax
Insurance
Capital recovery*
Total 1C
Total Annual Cost (rounded)
60% of sum of operating, supv., & maint. labor &
maint. materials = 0.6(25,920 + 3,888 +14, 256 +
14,256) =
2% of Total Capital Investment = 0.02($412,315)
1% of Total Capital Investment = 0.01($412,315)
1% of Total Capital Investment = 0.01(9412,315)
0.1175(412,315 - 2,809 - 13,220 x 1.08)
$25,920
3,888
14,256
14,256
9,845
48,323
8,294
148,114
$272,896
34,992
8,246
4,123
4,123
46,439
$97,923
$371,000
* The 1.08 factor is for freight and sales taxes.
6 The capital recovery cost factor, CRF, is a function of the fabric filter or equipment life
and the opportunity cost of the capital (i.e., interest rate). For example, for a 20 year
equipment life and a 10% interest rate, CRF = 0.1175.
5-49
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Pj = 100 psig
cleaning interval = 10 min
We further assume that a G/C of 4.69 ft/min is a good estimate of the
mean face velocity over the duration of the filtering cycle.
W0 = dV9
= 4-£ x „ *ib x 4.69 ?- x 10 min
ft3 7,000 gr nun
= 0.0268 lb/ft2
ft , , n«s in. H20/(ft/min)
AP = 6.08 x 4.69-V x (100 psig)'0'65 + 15 ' ';. -
nun ID/it
Ib ft
x0.0268-;-5 x 4.69 —
ft mm
' = 3.32 in. H20 across the fabric (when fully loaded).
We will assume that the baghouse structure and the ductwork contribute
an additional 3 in. H20 and 4 in. H20, respectively. The total pressure
drop is, therefore, 10.3 inches.
The total annual cost is $371,000, nearly half of which is for ash disposal.
If a market for the fly ash could be found, the total annual cost would be
greatly reduced. For example, if $2/ton were received for the ash, the total
annual cost would drop to $208,000 ($371,000 - $148,000 - $14,800), or
56% of the cost when no market exists. Clearly, the total annual cost is
extremely sensitive to the value chosen for the dust disposal cost in this
case. In this and in similar cases, this value should be selected with care.
As discussed in the Design Procedures Section (Section 5.2), an elec-
trostatically enhanced baghouse (an emerging technology) may have up to
30% lower total annual costs than the baghouse estimated in the example
problem.
5-50
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-------
5.5 Acknowledgment
We gratefully acknowledge the following companies for contributing data
to this section:
• Aget Manufacturing Company (Adrian, Michigan)
• BACT Engineering, Inc. (Arlington Heights, Illinois)
• The BHA Group (Kansas City, Missouri)
• Dustex Corporation (Charlotte, North Carolina)
• Fuller Company (Bethlehem, Pennsylvania)
• W. E. Gore and Associates, Inc. (Elkton, Maryland)
• Griffin Environmental Company, Inc. (Syracuse, New York)
• W. W. Sly Manufacturing Company (Cleveland, Ohio)
• Zurn Industries, Inc. (Birmingham, Alabama)
5-51
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References
[1] Van Osdell, D. W., M. B. Ranade, G. P. Greiner, and D. F. Furlong,
Electrostatic Augmentation of Fabric Filtration: Pulse-Jet Pilot Unit
Experience, November 1982 (EPA-600/7-82-062).
[2] Viner, A. S., G. P. Greiner, D. F. Furlong, and R. G. Hurst, Pilot-Scale
Evaluation of Top-Inlet and Advanced Electrostatic Filtration, October
1986 (EPA-600/7-86-042).
[3] Donovan, R. P., Fabric Filtration For Combustion Sources, Marcel
Dekker, Inc., New York, 1985.
[4] Turner, J. H., "Bag Filtration," in Handbook of Multiphase Systems,
ed. by G. Hetsroni, Hemisphere, New York, 1982.
[5] Turner, J. H., and J. D. McKenna, "Control of Particles by Filters,"
in Handbook of Air Pollution Technology, ed. by S. Calvert and H.
Englund, John Wiley & Sons, New York, 1984.
[6] Penny, G. W., Electrostatic Effects in Fabric Filtration: Volume I.
Fields, Fabrics, and Particles (Annotated Data), September 1978
(EPA-600/7-78-142A[NTIS PB 288576]).
[7] Frederick, E. R., Electrostatic Effects in Fabric Filtration: Volume II.
Triboelectric Measurements and Bag Performance, July 1978 (EPA-
600/7-78-lA2B[NTIS PB 287207]).
[8] Frederick, E. R., Electrical Effects in Particulate Matter Processes,
Filter Media Specification, Pittsburgh, 1987.
[9] Dennis, R., and H. A. Klemm, "Modeling Concepts for Pulse Jet Fil-
tration." JAPCA, 30(1), January 1980.
5-52
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[10] Leith, D. and M. J. Ellenbecker, "Theory for Pressure Drop in a Pulse-
Jet Cleaned Fabric Filter." Atm. Environment, 14, 1980, pp. 845-852.
[11] Koehler, J. L. and D. Leith, "Model Calibration for Pressure Drop in
a Pulse-Jet Cleaned Fabric Filter," Atm. Environment, 17(10), 1983,
pp. 1909-1913.
[12] Northrop Services, Inc. Fabric Filter Workshop Reference Materials,
1977 Workshop, Air Pollution Training Institute.
[13] Viner, A. S., and B. R. Locke, Cost and Performance Models for Elec-
trostatically Stimulated Fabric Filters, April 1984 (EPA 600/8-84-016).
[14] Vatavuk, W. M., and R. B. Neveril, "Estimating Costs of Air-Pollution
Control Systems, Part XI: Estimate the Size and Cost of Baghouses,"
Chemical Engineering, March 22, 1982, pp. 153-158.
[15] Frey, R. F., and T. V. Reinauer, "New Filter Rate Guide," Air Engi-
neering, 30 April 1964.
[16] Dennis, R., et al., Filtration Model for Coal Fly Ash with Glass Fabrics,
August 1977 (EPA-600/7-77-084 [NTIS PB 276489]).
[17] Owen, M. K. and A. S. Viner, Microcomputer Programs for Particulate
Control, June 1985 (EPA-600/8-85-025a).
[18] Dennis, R. and H. A. Klemm, Fabric Filter Model Change: Vol. I,
Detailed Technical Report, February 1979 (EPA-600/7-79-043a [NTIS
PB 293551]).
[19] Viner, A. S., et al., "Comparison of Baghouse Test Results with the
GCA/EPA Design Model", JAPCA, 34(8), August 1984.
[20] Reigel, S. A. and R. P. Bundy. "Why the Swing to Baghouses?"%Power,
121-1, January 1977, pp. 68-73.
[21] ETS, Inc., Roanoke, VA.
[22] Fuller Company, Bethlehem, PA.
[23] R. S. Means Company, Inc., Means Square Foot Costs 1986, Kingston,
MA.
[24] Vatavuk, W. M., and R. B. Neveril, "Estimating Costs of Air-Pollution
Control Systems, Part II: Factors for Estimating Capital and Operat-
ing Costs," Chemical Engineering, November 3, 1980, pp. 157-162.
5-53
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[25] Personal communication from Frank Smith, Griffin Environmental, to
Jim Turner, Research Triangle Institute, November 8, 1988.
[26] Perry, Robert H., ei a/., Perry's Chemical Engineers' Handbook
(Fourth Edition), McGraw-Hill, New York, 1963, p. 6-20.
[27] Personal communication from Gary Greiner, ETS, Inc., to Jim Turner,
Research Triangle Institute, October 24, 1986.
[28] Perry, Robert H., et a/., Perry's Chemical Engineers' Handbook (Sixth
Edition), McGraw-Hill, New York, 1984.
[29] McKenna, J. D., J. H. Turner, D. Furlong, and D. S. Beachler, Fabric
Filters-Baghouses, I, Theory, Design and Selection, in preparation,
ETS, Inc., Roanoke, VA.
5-54
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Chapter 6
ELECTROSTATIC
PRECIPITATORS
James H. Turner
Phil A. Lawless
Toshiaki Yamamoto
David W. Coy
Research Triangle Institute
Research Triangle Park, NC 27709
Gary P. Greiner
John D. McKenna
ETS, Inc.
Roanoke, VA 24018-4394
William M. Vatavuk
Standards Development Branch, OAQPS
U. S. Environmental Protection Agency
Research Triangle Park, NC 27711
November, 1989
6-1
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Contents
6.1 Process Description 6-4
6.1.1 Introduction 6-4
6.1.2 Types of ESPs 6-5
6.1.2.1 Plate-Wire Precipitators 6-5
6.1.2.2 Flat Plate Precipitators 6-10
6.1.2.3 Tubular Precipitators 6-10
6.1.2.4 Wet Precipitators 6-11
6.1.2.5 Two-Stage Precipitators 6-11
6.1.3 Auxiliary Equipment 6-12
6.1.4 Electrostatic Precipitation Theory 6-14
6.1.4.1 Electrical Operating Point 6-15
6.1.4.2 Particle Charging 6-17
.6.1.4.3 Particle Collection 6-19
6.1.4.4 Sneakage and Rapping Reentrainment .... 6-21
6.2 ESP Design Procedure 6-23
6.2.1 Specific Collecting Area 6-23
6.2.1.1 SCA Procedure with Known Migration Veloc-
ity 6-24
6.2.1.2 Full SCA Procedure 6-26
6.2.1.3 Specific Collecting Area for Tubular Precipi-
tators 6-34
6-2
-------
6.2.2 Flow Velocity 6-34
6.2.3 Pressure Drop Calculations 6-36
6.2.4 Particle Characteristics 6-37
6.2.5 Gas Characteristics 6-39
6.2.6 Cleaning 6-40
6.2.7 Construction Features 6-40
6.3 Estimating Total Capital Investment 6-42
6.3,1 Equipment Cost 6-42
6.3.1.1 ESP Costs 6-42
6.3.1.2 Retrofit Cost Factor 6-46
6.3.1.3 Auxiliary Equipment 6-47
6.3.1.4 Costs for Two-Stage Precipitators 6-47
6.3.2 Total Purchased Cost 6-49
6.3.3 Total Capital Investment (TCI) 6-50
6.4 Estimating Total Annual Costs 6-50
6.4.1 Direct Annual Costs 6-50
6.4.1.1 Operating and Supervisory Labor 6-52
6.4.1.2 Operating Materials 6-53
6.4.1.3 Maintenance 6-53
6.4.1.4 Electricity 6-53
6.4.1.5 Fuel 6-55
6.4.1.6 Water 6-55
6-3
-------
6.4.1.7 Compressed Air 6-55
6.4.1.8 Dust Disposal 6-55
6.4.1.9 Wastewater Treatment 6-56
6.4.1.10 Conditioning Costs 6-56
6.4.2 Indirect Annual Costs 6-56
6.4.3 Recovery Credits 6-57
6.4.4 Total Annual Cost 6-57
6.4.5 Example Problem 6-58
6.4.5.1 Design SCA 6-58
6.4.5.2 ESP Cost 6-62
6.4.5.3 Costs of Auxiliaries 6-62
6.4.5.4 Total Capital Investment 6-63
6.4.5.5 Annual Costs-Pressure Drop 6-63
6.4.5.6 Total Annual Cost 6-63
6.5 Acknowledgments 6-66
Appendix 6A - Effects of Material Thickness and Type on ESP Costs 6-67
References 6-73
6.1 Process Description
6.1.1 Introduction
An electrostatic precipitator (ESP) is a particle control device that uses
electrical forces to move the particles out of the flowing gas stream and
6-4
-------
onto collector plates. The particles are given an electrical charge by forcing
them to pass through a corona, a region in which gaseous ions flow. The
electrical field that forces the charged particles to the walls comes from
electrodes maintained at high voltage in the center of the flow lane.
Once the particles are collected on the plates, they must be removed
from the plates without reentraining them into the gas stream. This is
usually accomplished by knocking them loose from the plates, allowing the
collected layer of particles to slide down into a hopper from which they
are evacuated. Some precipitators remove the particles by intermittent or
continuous washing with water.
Figure 6.1 is an illustration of an ESP with its various components
identified. Figure 6.2 shows two variations of charging electrode/collector
electrode arrangements used in ESPs.
6.1.2 Types of ESPs
ESPs are configured in several ways. Some of these configurations have
been developed for special control action, and others have evolved for eco-
nomic reasons. The types that will be described here are (1) the plate-wire
precipitator, the most common variety; (2) the flat plate precipitator, (3)
the tubular precipitator; (4) the wet precipitator, which may have any of
the previous mechanical configurations; and (5) the two-stage precipitator.
6.1.2.1 Plate-Wire Precipitators
Plate-wire ESPs are used in a wide variety of industrial applications, in-
cluding coal-fired boilers, cement kilns, solid waste incinerators, paper mill
recovery boilers, petroleum refining catalytic cracking units, sinter plants,
basic oxygen furnaces, open hearth furnaces, electric arc furnaces, coke oven
batteries, and glass furnaces.
In a plate-wire ESP, gas flows between parallel plates of sheet metal
and high-voltage electrodes. These electrodes are long wires weighted and
hanging between the plates or are supported there by mast-like structures
(rigid frames). Within each flow path, gas flow must pass each wire in
sequence as it flows through the unit.
6-5
-------
VOtTAOC *Uf9Q*T HMUIATOM
Figure 6.1: Electrostatic Precipitator Components
(Courtesy of the Industrial Gas Cleaning Institute)
6-6
-------
CMUCKMPllll
• -Mil
Plan vww ol McU4 tP eteciioOe* «MII
T
COUKMMfUH
Plan vww o< Convamunal EP etocuodo* won lypical clmiaiuiioin
Figure 6.2: Flat-plate and Plate-wire ESP Configurations
(Courtesy of United McGill Corporation)
6-7
-------
The plate-wire ESP allows many flow lanes to operate in parallel, and
each lane can be quite tall. As a result, this type of precipitator is well
suited for handling large volumes of gas. The need for rapping the plates
to dislodge the collected material has caused the plate to be divided into
sections, often three or four in series with one another, which can be rapped
independently. The power supplies are often sectionalized in the same way
to obtain higher operating voltages, and further electrical sectionalization
may be used for increased reliability. Dust also deposits on the discharge
electrode wires and must be periodically removed similarly to the collector
plate.
The power supplies for the ESP convert the industrial ac voltage (220 to
480 V) to pulsating dc voltage in the range of 20,000 to 100,000 V as needed.
The supply consists of a step-up transformer, high-voltage rectifiers, and
sometimes filter capacitors. The unit may supply either half-wave or full-
wave rectified dc voltage. There are auxiliary components and controls
to allow the voltage to be adjusted to the highest level possible without
excessive sparking and to protect the supply and electrodes in the event a
heavy arc or short-circuit occurs.
The voltage applied to the electrodes causes the air between the elec-
trodes to break down electrically, an action known as a "corona". The
electrodes usually are given a negative polarity because a negative corona
supports a higher voltage than a positive corona before sparking occurs.
The ions generated in the corona follow electric field lines from the wires
to the collecting plates. Therefore, each wire establishes a charging zone
through which the particles must pass.
Particles passing through the charging zone intercept some of the ions,
which become attached. Small aerosol particles (<1 fj,m diameter) can
absorb tens of ions before their total charge becomes large enough to repel
further ions, and large particles (>10 pm diameter) can absorb tens of
thousands. The electrical forces are therefore much stronger on the large
particles.
As the particles pass each successive wire, they are driven closer and
closer to the collecting walls. The turbulence in the gas, however, tends to
keep them uniformly mixed with the gas. The collection process is therefore
a competition between the electrical and dispersive forces. Eventually, the
particles approach close enough to the walls so that the turbulence drops
to low levels and the particles are collected.
6-8
-------
If the collected particles could be dislodged into the hopper without
losses, the ESP would be extremely efficient. The rapping that dislodges the
accumulated layer also projects some of the particles (typically 12 percent
for coal fly ash) back into the gas stream. These reentrained particles are
then processed again by later sections, but the particles reentrained in the
last section of the ESP have no chance to be recaptured and so escape the
unit.
Practical considerations of passing the high voltage into the space be-
tween the lanes and allowing for some clearance above the hoppers to sup-
port and align electrodes leave room for part of the gas to flow around the
charging zones. This is called "sneakage" and amounts to 5 to 10 percent of
the total flow. Antisneakage baffles usually are placed to force the sneakage
flow to mix with the main gas stream for collection in later sections. But,
again, the sneakage flow around the last section has no opportunity to be
collected.
These losses play a significant role in the overall performance of an ESP.
Another major factor is the resistivity of the collected material. Because
the particles form a continuous layer on the ESP plates, all the ion current
must pass through the layer to reach the ground plates. This current creates
an electric field in the layer, and it can become large enough to cause local
electrical breakdown. When this occurs, new ions of the wrong polarity
are injected into the wire-plate gap where they reduce the charge on the
particles and may cause sparking. This breakdown condition is called "back
corona".
Back corona is prevalent when the resistivity of the layer is high, usually
above 2 x 10n ohm-cm. For lower resistivities, the operation of the ESP
is not impaired by back corona, but resistivities much higher than 2 x 1011
ohm-cm considerably reduce the collection ability of the unit because the
severe back corona causes difficulties in charging the particles. At resistiv-
ities below 108 ohm-cm, the particles are held on the plates so loosely that
rapping and nonrapping reentrainment become much more severe. Care
must be taken in measuring or estimating resistivity because it is strongly
affected by variables such as temperature, moisture, gas composition, par-
ticle composition, and surface characteristics.
6-9
-------
6.1.2.2 Flat Plate Precipitators
A significant number of smaller precipitators (100,000 to 200,000 acfm) use
flat plates instead of wires for the high-voltage electrodes. The flat plates
(United McGill Corporation patents) increase the average electric field that
can be used to collect the particles, and they provide an increased surface
area for the collection of particles. Corona cannot be generated on flat
plates by themselves, so corona-generating electrodes are placed ahead of
and sometimes behind the flat plate collecting zones. These electrodes may
be sharp-pointed needles attached to the edges of the plates or indepen-
dent corona wires. Unlike plate-wire or tubular ESPs, this design operates
equally well with either negative or positive polarity. The manufacturer
has chosen to use positive polarity to reduce ozone generation.
A flat plate ESP operates with little or no corona current flowing through
the collected dust, except directly under the corona needles or wires. This
has two consequences. The first is that the unit is somewhat less suscepti-
ble to back corona than conventional units are because no back corona is
generated in the collected dust, and particles charged with both polarities
of ions have large collection surfaces available. The second consequence is
that the lack of current in the collected layer causes an electrical force that
tends to remove the layer from the collecting surface; this can lead to high
rapping losses.
•
Flat plate ESPs seem to have wide application for high-resistivity par-
ticles with small (1 to 2 /zm) mass median diameters (MMDs). These
applications especially emphasize the strengths of the design because the
electrical dislodging forces are weaker for small particles than for large ones.
Fly ash has been successfully collected with this type of ESP, but low-flow
velocity appears to be critical for avoiding high rapping losses.
6.1.2.3 Tubular Precipitators
The original ESPs were tubular like the smokestacks they were placed on,
with the high-voltage electrode running along the axis of the tube. Tubular
precipitators have typical applications in sulfuric acid plants, coke oven
by-product gas cleaning (tar removal), and, recently, iron and steel sinter
plants. Such tubular units are still used for some applications, with many
tubes operating in parallel to handle increased gas flows. The tubes may
6-10
-------
be formed as a circular, square, or hexagonal honeycomb with gas flowing
upwards or downwards. The length of the tubes can be selected to fit
conditions. A tubular ESP can be tightly sealed to prevent leaks of material,
especially valuable or hazardous material.
A tubular ESP is essentially a one-stage unit and is unique in having
all the gas pass through the electrode region. The high-voltage electrode
operates at one voltage for the entire length of the tube, and the current
varies along the length as the particles are removed from the system. No
sneakage paths are around the collecting region, but corona nonuniformities
may allow some particles to avoid charging for a considerable fraction of
the tube length.
Tubular ESPs comprise only a small portion of the ESP population and
are most commonly applied where the particulate is either wet or sticky.
These ESPs, usually cleaned with water, have reentrainment losses of a
lower magnitude than do the dry particulate precipitators.
6.1.2.4 Wet Precipitators
Any of the precipitator configurations discussed above may be operated
with wet walls instead of dry. The water flow may be applied intermit-
tently or continuously to wash the collected particles into a sump for dis-
posal. The advantage of the wet wall precipitator is that it has no problems
with rapping reentrainment or with back corona. The disadvantage is the
increased complexity of the wash and the fact that the collected slurry
must be handled more carefully than a dry product, adding to -the expense
of disposal.
6.1.2.5 Two-Stage Precipitators
The previously described precipitators are all parallel in nature; i.e., the
discharge and collecting electrodes are side by side. The two-stage precip-
itator invented by Penney is a series device with the discharge electrode,
or ionizer, preceding the collector electrodes. For indoor applications, the
unit is operated with positive polarity to limit ozone generation.
Advantages of this configuration include more time for particle charging,
6-11
-------
Spray Cooler
Hood
Oinct Exhautt
Seraw Conwyor
Figure 6.3: Control Device and Typical Auxilary Equipment
less propensity for back corona, and economical construction for small sizes.
This type of precipitator is generally used for gas flow volumes of 50,000
acfm and less and is applied to submicrometer sources emitting oil mists,
smokes, fumes, or other sticky particulates because there is little electrical
force to hold the collected particulates on the plates. Modules consisting of
a mechanical prefilter, ionizer, collecting-plate cell, after-filter, and power
pack may be placed in parallel or series-parallel arrangements. Precon-
ditioning of gases is normally part of the system. Cleaning may be by
water wash of modules removed from the system up to automatic, in-place
detergent spraying of the collector followed by air-blow drying.
Two-stage precipitators are considered to be separate and distinct types
of devices compared to large, high-gas-volume, single-stage ESPs. The
smaller devices are usually sold as pre-engineered, package systems.
6.1.3 Auxiliary Equipment
Typical auxiliary equipment associated with an ESP system is shown sche-
matically in Figure 6.3. Along with the ESP itself, a control system usually
includes the following auxiliary equipment: a capture device (»'. c., hood or
6-12
-------
direct exhaust connection); ductwork; dust removal equipment (screw con-
veyor, etc.); fans, motors, and starters; and a stack. In addition, spray
coolers and mechanical collectors may be needed to precondition the gas
before it reaches the ESP. Capture devices are usually hoods that exhaust
pollutants into the ductwork or are direct exhaust couplings attached to a
combustor or process equipment. These devices are usually refractory lined,
water cooled, or simply fabricated from carbon steel, depending on the gas-
stream temperatures. Refractory or water-cooled capture devices are used
where the wall temperatures exceed 800° F; carbon steel is used for lower
temperatures. The ducting, like the capture device, should be water cooled,
refractory, or stainless steel for hot processes and carbon steel for gas tem-
peratures below approximately 1,150°F (duct wall temperatures <800°F).
The ducts should be sized for a gas velocity of approximately 4,000 ft/min
for the average case to prevent particle deposition in the ducts. Large or
dense particles might require higher velocities, but rarely would lower ve-
locities be used. Spray chambers may be required for processes where the
addition of moisture, or decreased temperature or gas volume, will improve
precipitation or protect the ESP from warpage. For combustion processes
with exhaust gas temperatures below approximately 700°F, cooling would
not be required, and the exhaust gases can be delivered directly to the
precipitator.
When much of the pollutant loading consists of relatively large particles, :
mechanical collectors, such as cyclones, may be used to reduce the load on
the ESP, especially at high inlet concentrations. The fans provide the
motive power for air movement and can be mounted before or after the
ESP. A stack, normally used, vents the cleaned stream to the atmosphere.
Screw conveyors or pneumatic systems are often used to remove captured
dust from the bottom of the hoppers.
Wet ESPs require a source of wash water to be injected or sprayed near
the top of the collector plates either continuously or at timed intervals. The
water flows with the collected particles into a sump from which the fluid is
pumped, A portion of the fluid may be recycled to reduce the total amount
of water required. The remainder is pumped directly to a settling pond or
passed through a dewatering stage, with subsequent disposal of the sludge.
Gas conditioning equipment to improve ESP performance by changing
dust resistivity is occasionally used as part of the original design, but more
frequently it is used to upgrade existing ESPs. The equipment injects
6-13
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an agent into the gas stream ahead of the ESP. Usually, the agent mixes
with the particles and alters their resistivity to promote higher migration
velocity, and thus higher collection efficiency. However, electrical properties
of the gas may change, rather than dust resistivity. For instance, cooling the
gas will allow more voltage to be applied before sparking occurs. Significant
conditioning agents that are used include SOa, H2S04, sodium compounds,"
^mmomaTand Water, but the major conditioning agent by usage is SO3. A
typical dosVrate ior any orihe gaseous~agents is 10 to 30'ppm by~volume.
The equipment required for conditioning depends on the agent being
used. A typical SOa conditioner requires a supply of molten sulfur. It is
stored in a heated vessel and supplied to a burner, where it is oxidized to
SO2. The S02 gas is passed over a catalyst for further oxidation to S03. The
SO3 gas is then injected into the flue gas stream through a multi-outlet set
of probes that breach a duct. In place of a sulfur burner to provide S02,
liquid SO2 may be vaporized from a storage tank. Although their total
annual costs are higher, the liquid SO2 systems have lower capital costs
and are easier to operate than the molten sulfur based systems.
Water or ammonia injection requires a set of spray nozzles in the duct,
along with pumping and control equipment.
Sodium conditioning is often done by coating the coal on a conveyor
with a powder compound or a water solution of the desired compound.
A hopper or storage tank is often positioned over the conveyor for this
purpose.
6.1.4 Electrostatic Precipitation Theory
The theory of ESP operation requires many scientific disciplines to describe
it thoroughly. The ESP is basically an electrical machine. The principal
actions are the charging of particles and forcing them to the collector plates.
The amount of charged particulate matter affects the electrical operating
point of the ESP. The transport of the particles is affected by the level of
turbulence in the gas. The losses mentioned earlier, sneakage and rapping
reentrainment, are major influences on the total performance of the system.
The particle properties also have a major effect on the operation of the unit.
The following subsections will explain the theory behind (1) electrical
6-14
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operating points in the ESP, (2) particle charging, (3) particle collection,
and (4) sneakage and rapping reentrainment. General references for these
topics are White [1] or Lawless and Sparks [2].
6.1.4.1 Electrical Operating Point
The electrical operating point of an ESP section is the value of voltage and
current at which the section operates. As will become apparent, the best
collection occurs when the highest electric field is present, which roughly
corresponds to the highest voltage on the electrodes. In this work, the
term "section" represents one set of plates and electrodes in the direction
of flow. This unit is commonly called a "field", and a "section" or "bus
section" represents a subdivision of a "field" perpendicular to the direction
of flow. In an ESP model and in sizing applications, the two terms "section"
and "field" are used equivalently because the subdivision into bus sections
should have no effect on the model. This terminology has probably arisen
because of the frequent use of the word "field" to refer to the electric field.
The lowest acceptable voltage is the voltage required for the formation of
a corona, the electrical discharge that produces ions for charging particles.
The (negative) corona is produced when an occasional free electron near
the high-voltage electrode, produced by a cosmic ray, gains enough energy
from the electric field to ionize the gas and produce more free electrons.
The electric field for which this process is self-sustained has been determined
experimentally. For round wires, the field at the surface of the wire is given
by:
Ee = 3.126 x 10a(t[l + 0.0301(dp/rw)°-5] (6.1)
where
Ee = corona onset field at the wire surface (V/m)
dr = relative gas density, referred to 1 atm pressure
and 20° C (dimensionless)
rw = radius of the wire, meters (m)
This is the field required to produce "glow" corona, the form usually
seen in the laboratory on smooth, clean wires. The glow appears as a uni-
form, rapidly moving diffuse light around the electrode. After a period of
operation, the movement concentrates into small spots on the wire surface,
and the corona assumes a tuft-like appearance. The field required to pro-
duce "tuft" corona, the form found in full-scale ESPs, is 0.6 times the value
6-15
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ofEc.
The voltage that must be applied to the wire to obtain this value of field,
Vc, is found by integrating the electric field from the wire to the plate. The
field follows a simple "1/r" dependence in cylindrical geometry. This leads
to a logarithmic dependence of voltage on electrode dimensions. In the
plate-wire geometry, the field dependence is somewhat more complex, but
the voltage still shows the logarithmic dependence. Vc is given by:
Ve = EcrwlJf\ (6.2)
\TW/
where
Vc = corona onset voltage (V)
i _ f outer cylinder radius for tubular ESP (m)
\ 4/7T x (wire-plate separation) for plate-wire ESP (m)
No current will flow until the voltage reaches this value, but the amount
of current will increase steeply for voltages above this value. The maximum
current density (amperes/square meter) on the plate or cylinder directly
under the wire is given by:
V2
J = Vfj3 (6-3)
where
j — maximum current density (A/m2)
fj, = ion mobility (m2/Vs) (meter2/volt second)
e = free space permittivity (8.845 x 10~12 F/m)(Farad/meter)
V — applied voltage (V)
L = shortest distance from wire to collecting surface (m)
For tuft corona, the current density is zero until the corona onset voltage
is reached, when it jumps almost to this value of j within a few hundred
volts, directly under a tuft.
The region near the wire is strongly influenced by the presence of ions
there, and the corona onset voltage magnitude shows strong spatial varia-
tions. Outside the corona region, it is quite uniform.
The electric field is strongest along the line from wire to plate and is
approximated very well, except near the wire, by:
Emax = V/L (6.4)
6-16
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where
Emaz = maximum field strength (V/m)
When the electric field throughout the gap between the wire and the plate
becomes strong enough, a spark will occur, and the voltage cannot be
increased without severe sparking occurring. The field at which sparking
occurs is not sharply defined, but a reasonable value is given by:
1>9S
E. = 6.3 x 10s (Pj (6.5)
where
E, = sparking field strength (V/m)
T = absolute temperature (K)
P = gas pressure (atm)
This field would be reached at a voltage of, for example, 35,000 V for a
plate-wire spacing of 11.4 cm (4.5 in.) at a temperature of 149°C (300°F).
The ESP will generally operate near this voltage in the absence of back
corona. EmaB will be equal to or less than Ef .
Instead of sparking, back corona may occur if the electric field in the
dust layer, resulting from the current flow in the layer, reaches a critical
value of about 1 x 106 V/m. Depending on conditions, the back corona
may enhance"' sparking or may generate so much current that the voltage
cannot be raised any higher. The field in the layer is given by:
EI = 3P (6.6)
where
EI = electric field in dust layer (V/m)
p = resistivity of the collected material (ohm-m)
6.1.4.2 Particle Charging
Charging of particles takes place when ions bombard the surface of a par-
ticle. Once an ion is close to the particle, it is tightly bound because of the
image charge within the particle. The "image charge" is a representation of
the charge distortion that occurs when a real charge approaches a conduct-
ing surface. The distortion is equivalent to a charge of opposite magnitude
to the real charge, located as far below the surface as the real charge is
6-17
-------
above it. The notion of the fictitious charge is similar to the notion of an
image in a mirror, hence the name. As more ions accumulate on a particle,
the total charge tends to prevent further ionic bombardment.
There are two principal charging mechanisms: diffusion charging and
field charging. Diffusion charging results from the thermal kinetic energy of
the ions overcoming the repulsion of the ions already on the particle. Field
charging occurs when ions follow electric field lines until they terminate
on a particle. In general, both mechanisms are operative for all sizes of
particles. Field charging, however, adds a larger percentage of charge on
particles greater than about 2 fjun in diameter, and diffusion charging adds
a greater percentage on particles smaller than about 0.5
Diffusion charging, as derived by White [1], produces a logarithmically
increasing level of charge on particles, given by:
= ln(l + r) (6.7)
where
q(t) s= particle charge (C) as function of time, t, in seconds
r = particle radius (m)
k — Boltzmann's constant (j/K)
T = absolute temperature (K)
e = electron charge (1.67 x 10~19 C)
T = dimensionless time given by:
r =
kT
where
(6.8)
v = mean thermal speed of the ions (m/s)
N = ion number concentration near the particle (No./m3)
9 = real time (exposure time in the charging zone) (s)
Diffusion charging never reaches a limit, but it becomes very slow after
about three dimensionless time units. For fixed exposure times, the charge
on a particle is proportional to its radius.
Field charging also exhibits a characteristic time- dependence, given by:
q(t) = q.O/(9 + r'} (6.9)
6-18
-------
where
q, = saturation charge, charge at infinite time (C)
0 = real time (s)
T' = another dimensionless time unit
The saturation charge is given by:
q. = 12Trcr*E (6.10)
where
e = free space permittivity (F/m)
E = external electric field applied to the particle (V/m)
The saturation charge is proportional to the square of the radius, which
explains why field charging is the dominant mechanism for larger particles.
The field charging time constant is given by:
r' = te/Nefj. (6.11)
where
H = ion mobility
Strictly speaking, both diffusion and field charging mechanisms operate
at the same time on all particles, and neither mechanism is sufficient to
explain the charges measured on the particles. It has been found empirically
that a very good approximation to the measured charge is given by the
sum of the charges predicted by equations 6.7 and 6.9 independently of one
another:
*«(*) = **(*) + ?/(<) (6.12)
where
?tot(0 = particle charge due to both mechanisms
9<*(0 = particle charge due to diffusion charging
= particle charge due to field charging
6.1.4.3 Particle Collection
The electric field in the collecting zone produces a force on a particle pro-
portional to the magnitude of the field and to the charge:
Fe = qE (6.13)
6-19
-------
where
Fe = force due to electric field (N)
q = charge on particle (C)
E = electric field (V/m)
Because the field charging mechanism gives an ultimate charge propor-
tional to the electric field, the force on large particles is proportional to the
square of the field, which shows the advantage for maintaining as high a
field as possible.
The motion of the particles under the influence of the electric field is
opposed by the viscous drag of the gas. By equating the electric force and
the drag force component due to the electric field (according to Stokes'
law), we can obtain the particle velocity:
OTTT/r
where
v(q,E,r) = particle velocity (m/s)
q(E,r) = particle charge (C)
C(r) = Cunningham correction to Stokes' law (dimensionless)
TJ = gas viscosity (kg/ms)
The particle velocity is the rate at which the particle moves along the
electric field lines, i.e., toward the walls.
For a given electric field, this velocity is usually at a minimum for par-
ticles of about 0.5 pro. diameter. Smaller particles move faster because the
charge does not decrease very much, but the Cunningham factor increases
rapidly as radius decreases. Larger particles have a charge increasing as r2
and a viscous drag only increasing as r1; the velocity then increases as r.
Equation 6.14 gives the particle velocity with respect to still air. In the
ESP, the flow is usually very turbulent, with instantaneous gas velocities of
the same magnitude as the particles velocities, but in random directions.
The motion of particles toward the collecting plates is therefore a statistical
process, with an average component imparted by the electric field and a
fluctuating component from the gas turbulence.
This statistical motion leads to an exponential collection equation, given
6-20
-------
by:
N(r) = NQ(r) x «p(-t;(r)/t>0) (6.15)
where
N(r) = particle concentration of size r at the exit of the col-
lecting zone (No./m3)
^o(i") = particle concentration of size r at the entrance of the
zone (No./m3)
v(r) = size-dependent particle velocity (m/s)
VG = characteristic velocity of the ESP (m/s), given by:
ro = Q/A = 1/SCA (6.16)
where
Q = volume flow rate of the gas (m3/s)
A = plate area for the ESP collecting zone (m2)
SCA = specific collection area (A/Q) (s/m)
When this collection equation is averaged over all the particle sizes and
weighted according to the concentration of each size, the Deutsch equation
results, with the penetration (fraction of particles escaping) given by:
p = exp(-we x SCA) (6.17)
where
p = penetration (fraction)
u>e = effective migration velocity for the particle
ensemble (m/s)
The efficiency is given by:
Eff(%) = 100(1 - p) (6.18)
and is the number most often used to describe the performance of an ESP.
6.1.4.4 Sneakage and Rapping Reentrainment
Sneakage and rapping reentrainment are best considered on the basis of
the sections within an ESP._^Qgakage_occurs when a part of thega
bypasses the collection zone of a section. GenerallyT'the portion of gas that
6-21
-------
bypasses the zone is thoroughly mixed with the gas that passes through the
zone before all the gas enters the next section. This mixing cannot always
be assumed, and when sneakage paths exist around several sections, the
performance of the whole ESP is seriously affected. To describe the effects
of sneakage and rapping reentrainment mathematically, we first consider
sneakage by itself and then consider the effects of rapping as an average
over many rapping cycles.
On the assumption that the gas is well mixed between sections, the
penetration for each section can be expressed as:
p. = SN + [(1 - Sir) x Pe(Q')} (6.19)
where
pt — section's fractional penetration
SN — fraction of gas bypassing the section (sneakage)
Pc(Q') = fraction of particles penetrating the collection zone,
which is functionally dependent on Q', the gas volume
flow in the collection zone, reduced by the sneakage
(m3/s)
The penetration of the entire ESP is the product of the section pene-
trations. The sneakage sets a lower limit on the penetration of particles
through the section.
To calculate the effects of rapping, we first calculate the amount of
material captured on the plates of the section. The fraction of material
that was caught is given by:
m/m0 = 1 - p. = 1 - SN - [(I - SN) x pe(Q')} (6.20)
where
m/m0 = mass fraction collected from the gas stream
This material accumulates until the plates are rapped, whereupon most
of the material falls into the hopper for disposal, but a fraction of it is
reentrained and leaves the section. Experimental measurements have been
conducted on fly ash ESPs to evaluate the fraction reentrained, which av-
erages about 12 percent.
The average penetration for a section, including sneakage and rapping
reentrainments, is:
Pt = SN + [(1 - 5*) x Pc(Q')} + RR(l - 5jv)[l - Pc(Q')} (6.21)
6-22
-------
where
RR = fraction reentrained
This can be written in a more compact form as:
P. = LF + [(1 - LF) x Pc(C?')] (6.22)
by substituting LF (loss factor) for SN + RR(l — SN)> These formulas
can allow for variable amounts of sneakage and rapping reentrainment for
each section, but there is no experimental evidence to suggest that it is
necessary.
Fly ash precipitators analyzed in this way have an average Sff of 0.07
and an RR of 0.12. These values are the best available at this time, but some
wet ESPs, which presumably have no rapping losses, have shown SN values
of 0.05 or less. These values offer a means for estimating the performance of
ESPs whose actual characteristics are not known, but about which general
statements can be made. For instance, wet ESPs would be expected to
have RR = 0, as would ESPs collecting wet or sticky particles. Particulate
materials with a much smaller mass mean diameter, MMD, than fly ash
would be expected to have a lower RR factor because they are held more
tightly to the plates and each other. Sneakage factors are harder to account
for; unless special efforts have been made in the design to control sneakage,
the 0.07 value should be used.
6.2 ESP Design Procedure
6.2.1 Specific Collecting Area
Specific collecting area (SCA) is a parameter used to compare ESPs and
roughly estimate their collection efficiency. SCA is the total collector plate
area divided by gas volume flow rate and has the units of s/m or s/ft. Since
SCA is the ratio of A/Q, it is often expressed as m2/(m3/s) or ft2/kacfm,
where kacfm is thousand acfm. SCA is also one of the most important fac-
tors in determining the capital and several of the annual costs (for example,
maintenance and dust disposal costs) of the ESP because it determines the
size of the unit. Because of the various ways in which SCA can be ex-
pressed, Table 6.1 gives equivalent SCAs in the different units for what
would be considered a small, medium, and large SCA.
6-23
-------
Table 6.1: Small, Medium, and Large SCAs as Expressed by Various Units
Units
ft'/kacfm
s/m
s/ft
Small Medium Large
100 400 900
19.7 78.8 177
6 24 54
5.080 ftVkacfm = 1 (s/m).
The design procedure is based on the loss factor approach of Lawless
and Sparks [2] and considers a number of process parameters. It can be
calculated by hand, but it is most conveniently used with a spreadsheet
program. For many uses, tables of effective migration velocities can be
used to obtain the SCA required for a given efficiency. In the following
subsection, tables have been calculated using the design procedure for a
number of different particle sources and for differing levels of efficiency. If
a situation is encountered that is not covered in these tables, then the full
procedure that appears in the subsequent subsection should be used.
6.2.1.1 SCA Procedure with Known Migration Velocity
If the migration velocity is known, then equation 6.17 can be rearranged to
give the SCA:
(6.23)
A graphical solution to equation 6.23 is given in Figure 6.4. The mi-
gration velocities have been calculated for three main precipitator types:
plate- wire, flat plate, and wet wall ESPs of the plate- wire type. The follow-
ing three tables, keyed to design efficiency as an easily quantified variable,
summarize the migration velocities under various conditions:
• In Table 6.2, the migration velocities are given for a plate-wire ESP
with conditions of no back corona and severe back corona; tempera-
tures appropriate for each process have been assumed.
6-24
-------
"•*
,00 .U UH,
400 4<0 100
PW 1,000 ft'/mta
«00 fM
Figure 6.4: Chart for Finding SCA
6-25
-------
• In Table 6.3, the migration velocities calculated for a wet wall ESP
of the plate-wire type assume no back corona and no rapping reen-
trainment.
• In Table 6.4, the flat plate ESP migration velocities are given only
for no back corona conditions because they appear to be less affected
by high-resistivity dusts than the plate-wire types.
It is generally expected from experience that the migration velocity will
decrease with increasing efficiency. In Tables 6.2 through 6.4, however, the
migration velocities show some fluctuations. This is because the number
of sections must be increased as the efficiency increases, and the changing
sectionalization affects the overall migration velocity. This effect is par-
ticularly noticeable, for example, in Table 6.4" for glass plants. When the
migration velocities in the tables are used to obtain SCAs for the different
efficiencies in the tables, the SCAs will increase as the efficiency increases.
6.2.1.2 Full SCA Procedure
The full procedure for determining the SCA for large plate-wire, flat plate,
and (with restrictions) tubular dry ESPs is given here. This procedure does
not apply to the smaller, two-stage precipitators because these are packaged
modules generally sized and sold on the basis of the waste gas volumetric
flow rate. Nor does this procedure apply to determining the SCA for wet
ESPs. The full procedure consists of the 15 steps given below:
Step 1 - Determine the design efficiency, Eff(%). Efficiency is the most
commonly used term in the industry and is the reference value for guaran-
tees; however, if it has not been specified, it can be computed as follows:
Eff(%) = 100 x (1 - outlet load/inlet load)
Step 2 - Compute design penetration, p:
p = l- (Eff/100)
6-26
-------
Table 6.2: Plate-wire ESP Migration Velocities
(cm/s)'
Design Efficiency, %
Particle Source
Bituminous coal fly ash*
Sub-bituminous coal fly ash in
tangential-fired boiler*
Other coal6
Cement kilnc
Glass plant"*
Iron/steel sinter plant dust with
mechanical precollector*
Kraft-paper recovery boiler6
Incinerator fly ash"
Copper reverberatory furnace-'
Copper converter*
Copper roaster'1
Coke plant combustion stack*
(no BC)
(BC)
(no BC)
(BC)
(no BC)
(BC)
(no BC)
(BC)
(no BC)
(BC)
(no BC)
(BC)
(no BC)
(no BC)
(no BC)
(no BC)
(no BC)
(no BC)
95
12.6
3.1
17.0
4.9
9.7
2.9
1.5
0.6
1.6
0.5
6.8
2.2
2.6
15.3
6.2
5.5
6.2
1.2>
99
10.1
2.5
11.8
3.1
7.9
2.2
1.5
0.6
1.6
0.5
6.2
1.8
2.5
11.4
4.2
4.4
5.5
—
99.5
9.3
2.4
10.3
2.6
7.9
2.1
1.8
0.5
1.5
0.5
6.6
1.8
3.1
10.6
3.7
4.1
5.3
—
99.9
8.2
2.1
8.8
2.2
7.2
1.9
1.8
0.5
1.5
0.5
6.3
1.7
2.9
9.4
2.9
3.6
4.8
—
BC = Back corona.
*To convert cm/s to ft/s, multiply cm/s by 0.0328. Computational procedure uses SI units, to con-
vert cm/s to m/s, multiply cm/s by 0.01. Assumes same particle size as given in full computational
procedure.
'At 300°F. Depending on individual furnace/boiler conditions, chemical nature of the fly ash, and
availability of naturally occurring conditioning agents (e.g., moisture in the gas stream), migration
velocities may vary considerably from these values. Likely values are in the range from back corona
to no back corona.
'At 600°F. ''At 500°F. «At 250°F. >4SO to 570 °F.
»500 to 700 °F. *600 to 680 °F. '360 to 450 °F.
'Data available only for inlet concentrations in the range of 0.02 to 0.2 g/s ma and for efficiencies
less than 90 percent.
6-27
-------
Table 6.3: Wet Wall Plate-wire ESP Migration Velocities
(Nc back corona, cm/s)°
Design efficiency, %
Particle Source6 95 99 99.5 99.9~
Bituminous coal fly ash 31.4 33.0 33.8 24.9
Sub-bituminous coal fly ash in 40.0 42.7 44.1 31.4
tangential-fired boiler
Other coal 21.1 21.4 21.5 17.0
Cement kiln 6.4 5.6 5.0 5.7
Glass plant 4.6 4.5 4.3 3.8
Iron/steel sinter plant dust with 14.0 13.7 13.3 11.6
mechanical precollector
"To convert cm/s to ft/s, multiply cm/s by 0.0328. Computational procedure
uses SI units; to convert cm/s to m/s, multiply cm/s by 0.01. Assumes same
particle size as given in full computational procedure.
*All sources assumed at 200 *F.
6-28
-------
Table 8.4: Flat Plate ESP Migration Velocities"
(No back corona, cm/s)fc
Design Efficiency, %
Particle Source ~9599 99.5™~99.9
Bituminous coal fly ashc 13.2 15.1 18.6 16.0
Sub-bituminous coal fly ash in tangential- 28.6 18.2 21.2 17.7
fired boiler6
Other coalc 15.5 11.2 15.1 13.5
Cement kiln* 2.4 2.3 3.2 3.1
Glass plant' 1.8 1.9 2.6 2.6
Iron/steel sinter plant dust with mechanical 13.4 12.1 13.1 12.4
precollectorc
Kraft-paper recovery boiler0 5.0 4.7 6.1 5.8
Incinerator fly ash' 25.2 16.9 21.1 18.3
"Assumes same particle size as given in full computational procedure. These values
give the grounded collector plate SCA, from which the collector plate area is derived.
In flat plate ESPs, the discharge or high-voltage plate area is typically 40 percent
of the ground-plate area. The flat-plate manufacturer usually counts all the plate
area (collector plates plus discharge plates) in meeting an SCA specification, which
means that the velocities tabulated above must be divided by 1.4 to be used on the
manufacturer's basis.
kTo convert cm/s to ft/s, multiply cm/a by 0.0328. Computational procedure uses SI
units; to convert cm/s to m/s, multiply cm/s by 0.01.
"At 300° F.
''At 600°F.
•At 500° F.
250°F.
6-29
-------
Step 3 - Compute or obtain the operating temperature, Tk, K. Temper-
ature in Kelvin is required in the calculations which follow.
Step 4 - Determine whether severe back corona is present. Severe back
corona usually occurs for dust resistivities above 2 x 1011 ohm-cm. Its
presence will greatly increase the size of the ESP required to achieve a
certain efficiency.
Step 5 - Determine the MMD of the inlet particle distribution MMD<
(fim). If this is not known, assume a value from the following table:
Source MMDj (/xm)
Bituminous coal 16
Sub-bituminous coal,
tangential boiler 21
Sub-bituminous coal,
other boiler types 10 to 15
Cement kiln 2 to 5
Glass plant 1
Wood burning boiler 5
Sinter plant, 50
with mechanical precollector 6
Kraft Process Recovery 2
Incinerators 15 to 30
Copper reverberatory furnace 1
Copper converter 1
Coke plant combustion stack 1
Unknown 1
Step 6 - Assume values for sneakage, Sjy, and rapping reentrainment,
RR, from the following table:
ESP Type
Plate-wire 0.07
Wet wall 0.05
Flat plate 0.10
6-30
-------
ESP/Ash Type RR
Coal fly ash, or not known 0.124
Wet wall 0.0
Flat plate with gas velocity
>1.5 m/s (not glass or cement) 0.15
Glass or cement 0.10
Step 7 - Assume values for the most penetrating size, MMD,,, and rap-
ping puff size, MMDr:
MMDp = 2
MMDr = 5 f*m for ash with MMD,- > 5
MMDr = 3 fj.m for ash with MMDj < 5 /xm
where
MMD,, = the MMD of the size distribution emerging from
a very efficient collecting zone
MMDr = the MMD of the size distribution of rapped/
reentrained material.
Step 8 - Use or compute the following factors for pure air:
eO = 8.845xlO~12 free space permittivity (F/m)
77 =. 1.72xlO-5(Tk/273)°-71 gas viscosity (kg/ms)
EM = 630,000 (273/Tk)1'66 electric field at sparking
(V/m)
LF = S;v 4- RR(1 - S#) loss factor (dimensionless)
For plate-wire ESPs:
Eavg = .Efcj/1.75 average field with no back
corona
Eavg = 0.7 x.Ew/1.75 average field with severe back
corona
average
corona
6-31
-------
For flat plate ESPs:
Eavg = EM*. 5/6.3 average field, no back corona,
positive polarity
Eavg = 0.7 x#Mx 5/6.3 average field, severe back
corona, positive polarity
Step 9 - Assume the smallest number of sections for the ESP, n, such
that LFn < p. Suggested values of n are:
Eff(%) n
<96.5 2
<99 3
<99.8 4
<99.9 5
>99.9 6
These values are for an LF of 0.185, corresponding to a coal fly ash
precipitator. The values are approximate, but the best results are for the
smallest allowable n.
Step 10 - Compute the average section penetration, p,:
P. = P1/n
Step 11 - Compute the section collection penetration, pe:
P.-LF
Pc =
l-LF
If the value of n is too small, then this value will be negative and n must
be increased.
Step 12 - Compute the particle size change factors, D and MMDrp, which
are constants used for computing the change of particle size from section
to section:
D = p. = SN + Pc(l - Sir) + RR(1 - SN)(\ - Pe)
6-32
-------
MMDrp = RR(1- SN)(l-Pc)MMDr/D
Step 13 - Compute a table of particle sizes for sections 1 through n:
Section MMD
I MMDi = MMD;
2 MMD2 = {MMDtX SN + [(I - Pc) x MMDP + Pex MMD^
xpe}/D + MMDrp
3 MMD3 = {MMD2x S* + [(1 - pc) x MMDP + Pex MMD2]
xpc}/D + MMDrp
n MMDn = {MMDn_tx SN + [(I - Pc) x MMDP + Pex
MMDn_a] xpc}/D + MMDrp
Step 14 - Calculate the SCA for sections 1 through n, using MMDn, 77,6,
and pc:
= -(7;/e) x (1 - Sir) x In (p«)/(E2.,x MMDa x lO'8)
SCAn = -(17/6) x (1 - SAT) xln (Pe)/(E^x MMDn x 10'6)
where the factor 10 e converts micrometers to meters. Note that the only
quantity changing in these expressions is MMDB; therefore, the following
relation can be used:
SCAn+1 = SCAn x MMDn/MMDn+!
Step 15 - Calculate the total SCA and the English SCA, ESCA:
SCA (s/m) = "
6-33
-------
ESCA (ft2/kacfm) = 5.080 x SCA (s/m)
This sizing procedure works best for pc values less than the value of
LF, which means the smallest value of n. Any ESP model is sensitive to
the values of particle diameter and electric field. This one shows the same
sensitivity, but the expressions for electric field are based on theoretical
and experimental values. The SCA should not be strongly affected by
the number of sections chosen; if more sections are used, the SCA of each
section is reduced.
6.2.1.3 Specific Collecting Area for Tubular Precipitators
The procedure given above is suitable for large plate-wire or flat plate ESPs,
but must be used with restrictions for tubular ESPs. Values of Sjv = 0.015
and RR = 0 are assumed, and only one section is used.
Table 6.5 gives migration velocities that can be used with equation 6.23
to calculate SCAs for several tubular ESP applications.
6.2.2 Flow Velocity
A precipitator collecting a dry particulate material runs a risk of nonrapping
(continuous) reentrainment if the gas velocity becomes too large. This effect
is independent of SCA and has been learned through experience. For fly.
ash applications, the maximum acceptable velocity is about 1.5 m/s (5 ft/s)
for plate-wire ESPs and about 1 m/s (3 ft/s) for flat plate ESPs. For low
resistivity applications, design velocities of 3 ft/s or less are common to
avoid nonrapping reentrainment. The frontal area of the ESP (W x H),
i.e., the area normal to the direction of gas flow, must be chosen to keep
gas velocity low and to accommodate electrical requirements (e.g., wire-to-
plate spacing) while also ensuring that total plate area requirements are
met. This area can be configured in a variety of ways. The plates can
be short in height, long in the direction of flow, with several in parallel
(making the width narrow). Or, the plates can be tall in height, short in
the direction of flow, with many in parallel making the width large). After
selecting a configuration, the gas velocity can be obtained by dividing the
6-34
-------
Table 6.5: Tubular ESP Migration Velocities"
(cm/s)fc
Particle Source
Cement kiln
Glass plant
Kraft-paper
recovery boiler
Incinerator
15 fan MMD
Wet, at 200°F
MMD (fjtm)
1
2
5
10
20
Design Efficiency, %
(no BC)
(BC)
(no BC)
(BC)
(no BC)
(no BC)
90
2.2-5.4
1.1-2.7
1.4
0.7
4.7
40.8
3.2
6.4
16.1
32.2
64.5
95
2.1-5
.1
1.0-2.6
1.3
0.7
4.4
39.
3.1
6.2
15.4
30.8
61.6.
BC = Back corona
"These rates were calculated on the basis of:
Sjv = 0.015, RR = 0, one section only.
These are in agreement with operating tubular ESPs; exten-
sion of results to more than one section is not recommended.
*To convert cm/a to ft/s, multiply cm/a by 0.0328.
6-35
-------
volume flow rate, Q, by the frontal area of the ESP:
V"" = \VH (6>24)
where
Vga» — gas velocity (m/s)
W = width of ESP entrance (m)
H = height of ESP entrance (m)
When meeting the above restrictions, this value of velocity also ensures
that turbulence is not strongly developed, thereby assisting in the capture
of particles.
6.2.3 Pressure Drop Calculations
The pressure drop in an ESP is due to four main factors:
• Diffuser plate (usually present)—(perforated plate at the inlet)
• Transitions at the ESP inlet and outlet
• Collection plate baffles (stiffeners) or corrugations
• Drag of the flat collection plate
The total pressure drop is the sum of the individual pressure drops, but
any one of these sources may dominate all other contributions to the pres-
sure drop. Usually, the pressure drop is not a design-driving factor, but it
needs to be maintained at an acceptably low value. Table 6.6 gives typical
pressure drops for the four factors. The ESP pressure drop, usually less
than about 0.5 in. H2O, is much lower than for the associated collection
system and ductwork. With the conveying velocities used for dust collected
in ESPs, generally 4,000 ft/min or greater, system pressure drops are usu-
ally in the range of 2 to 10 in. H20, depending upon the ductwork length
and configuration as well as the type(s) of preconditioning devices(s) used
upstream.
The four main factors contributing to pressure drop are described briefly
below.
6-36
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Table 6.6: Components of ESP Pressure Drop
Typical Pressure Drop (in. H20)
Component
Diffuser
Inlet transition
Outlet transition
Baffles
Collection plates
Total
Low
0.010
0.07
0.007
0.0006
0.0003
0.09
High
0.09
0.14
0.015
0.123
0.008
0.38
The diffuser plate is used to equalize the gas flow across the face of the
ESP. It typically consists of a flat plate covered with round holes of 5 to 7
cm diameter (2 to 2.5 in.) having an open area of 50 to 65 percent of the
total. Pressure drop is strongly dependent on the percent open area, but
is almost independent of hole size.
The pressure drop due to gradual enlargement at the inlet is caused by
the combined effects of flow separation and wall friction and is dependent
on the shape of the enlargement. At the ESP exit, the pressure drop caused
by a short, well-streamlined gradual contraction is small.
Baffles are installed on collection plates to shield the collected dust from
the gas flow and to provide a stiffening effect to keep the plates aligned
parallel to one another. The pressure drop due to the baffles depends on
the number of baffles, their protrusion into the gas stream with respect to
electrode-to-plate distance, and the gas velocity in the ESP.
The pressure drop of the flat collection plates is due to friction of the gas
dragging along the flat surfaces and is so small compared to other factors
that it may usually be neglected in engineering problems.
6.2.4 Particle Characteristics
Several particle characteristics are important for particle collection. It is
generally assumed that the particles are spherical or spherical enough to
6-37
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be described by some equivalent spherical diameter. Highly irregular or
elongated particles may not behave in ways that can be easily described.
The first important characteristic is the mass of particles in the gas
stream, i.e., the particle loading. This quantity usually is determined by
placing a filter in the gas stream, collecting a known volume of gas, and
determining the weight gain of the filter. Because the ESP operates over a
wide range of loadings as a constant efficiency device, the inlet loading will
determine the outlet loading directly. If the loading becomes too high, the
operation of the ESP will be altered, usually for the worse.
The second characteristic is the size distribution of the particles, often
expressed as the cumulative mass less than a given particle size. The size
distribution describes how many particles of a given size there are, which
is important because ESP efficiency varies with particle size. In practical
terms, an ESP will collect all particles larger than 10 fj,m in diameter better
than ones smaller than 10 ^m. Only if most of the mass in the particles is
concentrated above 10 /zm would the actual size distribution above 10 ^m
be needed.
In lieu of cumulative mass distributions, the size distribution is often
described by log-normal parameters. That is, the size distribution appears
as a probabilistic normal curve if the logarithm of particle size used is the
abscissa. The two parameters needed to describe a log-normal distribu-
tion are the mass median (or mean) diameter and the geometric standard
deviation.
The MMD is the diameter for which one-half of the particulate mass
consists of smaller particles and one-half is larger (see the Procedure, Step
5, of Subsection 6.2.1.2). If the MMD of a distribution is larger than about
3 /mi, the ESP will collect all particles larger than the MMD at least as
well as a 3 /zm particle, representing one-half the mass in the inlet size
distribution.
The geometric standard deviation is the equivalent of the standard de-
viation of the normal distribution: It describes how broad the size distri-
bution is. The geometric standard deviation is computed as the ratio of
the diameter corresponding to 84 percent of the total cumulative mass to
the MMD; it is always a number greater than 1. A distribution with parti-
cles of all the same size (monodisperse) has a geometric standard deviation
of 1. Geometric standard deviations less than 2 represent rather narrow
6-38
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distributions. For combustion sources, the geometric standard deviations
range from 3 to 5 and are commonly in the 3.5 to 4.5 range.
A geometric standard deviation of 4 to 5, coupled with an MMD of less
than 5 ^m, means that there is a substantial amount of submicrometer
material. This situation may change the electrical conditions in an ESP by
the phenomenon known as "space charge quenching", which results in high
operating voltages but low currents. It is a sign of inadequate charging
and reduces the theoretical efficiency of the ESP. This condition must be
evaluated carefully to be sure of adequate design margins.
6.2.5 Gas Characteristics
The gas characteristics most needed for ESP design are the gas volume
flow and the gas temperature. The volume flow, multiplied by the design
SCA, gives the total plate area required for the ESP. If the volume flow
is known at one temperature, it may be estimated at other temperatures
by applying the ideal gas law. Temperature and volume uncertainties will
outweigh inaccuracies of applying the ideal gas law.
The temperature of the gas directly affects the gas viscosity, which
increases with temperature. Gas viscosity is affected to a lesser degree
by the gas composition, particularly the water vapor content. In lieu of
viscosity values for a particular gas composition, the viscosity for air may
be used. Viscosity enters the calculation of SCA directly, as seen in Step
14 of the design procedure (page 6-33).
The gas temperature and composition can have a strong effect on the
resistivity of the collected particulate material. Specifically, moisture and
acid gas components may be chemisorbed on the particles in a, sufficient
amount to lower the intrinsic resistivity dramatically (orders of magni-
tude). For other types of materials, there is almost no effect. Although
it is not possible to treat resistivity here, the designer should be aware of
the potential sensitivity of the size of the ESP to resistivity and the factors
influencing it.
The choice of power supplies' size (current capacity and voltage) to be
used with a particular application may be influenced by the gas characteris-
tics. Certain applications produce gas whose density may vary significantly
6-39
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from typical combustion sources (density variation may result from tem-
perature, pressure, and composition). Gas density affects corona starting
voltages and voltages at which sparking will occur.
6.2.6 Cleaning
Cleaning the collected materials from the plates often is accomplished in-
termittently or continuously by rapping the plates severely with automatic
hammers or pistons, usually along their top edges, except in the case of wet
ESPs that use water. Rapping dislodges the material, which then falls down
the length of the plate until it lands in a dust hopper. The dust character-
istics, rapping intensity, and rapping frequency determine how much of the
material is reentrained and how much reaches the hopper permanently.
For wet ESPs, consideration must be given to handling waste waters.
For simple systems with innocuous dusts, water with particles collected by
the ESP may be discharged from the ESP system to a solids-removing clari-
fier (either dedicated to the ESP or part of the plant wastewater treatment
system) and then to final disposal. More complex systems may require
skimming and sludge removal, clarification in dedicated equipment, pH ad-
justment, and/or treatment to remove dissolved solids. Spray water from
the ESP preconditioner may be treated separately from the water used to
flood the ESP collecting plates, so that the cleaner of the two treated wa-
ters may be returned to the ESP. Recirculation of treated water to the
ESP may approach 100 percent.
The hopper should be designed so that all the material in it slides to the
very bottom, where it can be evacuated periodically, as the hopper becomes
full. Dust is removed through a valve into a dust-handling system, such
as a pneumatic conveyor. Hoppers often are supplied with auxiliary heat
to prevent the formation of lumps or cakes and the subsequent blockage of
the dust handling system.
6.2.7 Construction Features
The use of the term "plate-wire geometry" may be somewhat misleading.
It could refer to three different types of discharge electrodes: weighted
6-40
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wires hung from a support structure at the top of the ESP, wire frames in
which wires are strung tautly in a rigid support frame, or rigid electrodes
constructed from a single piece of fabricated metal. In recent years, there
has been a trend toward using wire frames or rigid discharge electrodes
in place of weighted wire discharge electrodes (particularly in coal-fired
boiler applications). This trend has been stimulated by the user's desire
for increased ESP reliability. The wire frames and rigid electrodes are
less prone to failure by breakage and are readily cleaned by impulse-type
cleaning equipment.
Other differences in construction result from the choice of gas passage
(flow lane) width or discharge electrode to collecting electrode spacing.
Typically, discharge to collecting electrode spacing varies from 11 to 19 cm
(4.3 to 7.5 in.). Having a large spacing between discharge and collecting
electrodes allows higher electric fields to be used, which tends to improve
dust collection. To generate larger electric fields, however, power supplies
must produce higher operating voltages. Therefore, it is necessary to bal-
ance the cost savings achieved with larger electrode spacing against the
higher cost of power supplies that produce higher operating voltages.
Most ESPs are constructed of mild steel. ESP shells are constructed
typically of 3/16 or 1/4 in. mild steel plate. Collecting electrodes are gen-
erally fabricated from lighter gauge mild steel. A thickness of 18 gauge is
common, but it will vary with size and severity of application.
Wire discharge electrodes come in varied shapes from round to square
or barbed. A diameter of 2.5 mm (0.1 in.) is common for weighted wires,
but other shapes used have much larger effective diameters, e.g., 64 mm
(0.25 in.) square electrodes.
Stainless steel may be used for corrosive applications, but it is uncom-
mon except in wet ESPs. Stainless steel discharge electrodes have been
found to be prone to fatigue failure in dry ESPs with impact-type elec-
trode cleaning systems.[3]
Precipitators used to collect sulfuric acid mist in sulfuric acid plants
are constructed of steel, but the surfaces in contact with the acid mist are
lead-lined. Precipitators used on paper mill black liquor recovery boilers
are steam-jacketed. Of these two, recovery boilers have by far the. larger
number of ESP applications.
6-41
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Table 6.7: Standard Options for Basic Equipment
Option Cost adder (%)
1 - Inlet and outlet nozzles and diffuser plates 8 to 10
2 - Hopper auxiliaries/heaters, level detectors 8 to 10
3 - Weather enclosure and stair access 8 to 10
4 - Structural supports 5
5 - Insulation 8 to 10
Total options 1 to 5 1.37 to 1.45xBase
6.3 Estimating Total Capital Investment
Total capital investment (TCI) for an ESP system includes costs for the
ESP structure, the internals, rappers, power supply, auxiliary equipment,
and the usual direct and indirect costs associated with installing or erecting
new control equipment. These costs, in second-quarter 1987 dollars, are
described in the following subsections.
6.3.1 Equipment Cost
6.3.1.1 ESP Costs
Five types of ESPs are considered: plate-wire, flat plate, wet, tubular, and
two-stage. Basic costs for the first two are taken from Figure 6.5, which
gives the flange-to-flange, field-erected price based on required plate area
and a rigid electrode design. This plate area is calculated from the sizing
information given previously for the four types. Adjustments are made
for standard options listed in Table 6.7. Costs for wet/tubular ESPs are
discussed under Recent Trends, below, and costs for two-stage precipitators
are given in a later subsection.
The costs are based on a number of actual quotes. Least squares lines
have been fitted to the quotes, one for sizes between 50,000 and 1,000,000
ft2, and a second for sizes between 10,000 and 50,000 ft2. An equation
6-42
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103
10s
2 3 4 8 87891
££
14JJO
Rigid
nth All
2 3 4 S 8 7 8 91
Options
^4
HUM
t»
h-a^
Fl«ng»teRi
10,000
100,000
PtntAraa-ft2
1X 106
Figure 6.5: Dry-type ESP Flange-to-flange Purchase Price vs. Plate Area
6-43
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is given for each line. Extrapolation below 10,000 or above 1,000,000 ft2
should not be used. The reader should not be surprised if quotes are ob-
tained that differ from these curves by as much as ±25 percent. Significant
savings can be had by soliciting multiple quotes. All units include the ESP
casing, pyramidal hoppers, rigid electrodes and internal collecting plates,
transformer rectifier (TR) sets and microprocessor controls, rappers, and
stub supports (legs) for 4 feet clearance below the hopper discharges. The
lower curve is the basic unit without the standard options. The upper
curve includes all of the standard options (see Table 6.7) that are normally
utilized in a modern system. These options add approximately 45 percent
to the basic cost of the flange-to-flange hardware. Insulation costs are for
3 in. of field-installed glass fiber encased in a metal skin and applied on the
outside of all areas in contact with the exhaust gas stream. Insulation for
ductwork, fan casings, and stacks must be calculated separately.
Impact of alternative electrode designs All three designs—rigid elec-
trode, weighted wire, and rigid frame—can be employed in most applica-
tions. Any cost differential between designs will depend on the combination
of vendor experience and site-specific factors that dictate equipment size
factors. The rigid frame design will cost up to 25 percent more if the mast
or plate height is restricted to the same used in other designs. Several ven-
dors can now provide rigid frame collectors with longer plates, and thus the
cost differential can approach zero.
The weighted wire design uses narrower plate spacings and more inter-
nal discharge electrodes. This design is being employed less; therefore, its
cost is increasing and currently is approximately the same as that for the
rigid electrode collector. Below about 15,000 ft2 of plate area, ESPs are of
different design and are not normally field erected, and the costs will be
significantly different from values extrapolated from Figure 6.5.
Impact of materials of construction: Metal thickness and stain-
less steel Corrosive or other adverse operating conditions may suggest
the specification of thicker metal sections in the precipitator. Reasonable
increases in metal sections result in minimal cost increases. For example,
collecting plates are typically constructed of 18 gauge mild steel. Most ESP
manufacturers can increase the section thickness by 25 percent without sig-
nificant design changes or increases in manufacturing costs of more than a
6-44
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few percent.
Changes in type of material can increase purchase cost of the ESP from
about 30 to 50 percent for type 304 stainless steel collector plates and
precipitator walls, and up to several hundred percent for more expensive
materials used for all elements of the ESP. Based on the type 304 stain-
less steel cost, the approximate factors given below can be used for other
materials:
Material Factor Reference(s)
Stainless steel, 316
Carpenter 20 CB-3
Monel-400
Nickel-200
Titanium
1.3
1.9
2.3
3.2
4.5
[4,5,6]
[6]
[4,6]
[6]
[6]
Appendix 6A provides more detail on the effects of material thickness
and type.
Recent trends Most of today's market (1987) is in the 50,000 to 200,000
ft2 plate area size range. ESP selling prices have increased very little over
the past 10 years because of more effective designs, increased competition
from European suppliers, and a shrinking utility market.
Design improvements have allowed wider plate spacings that reduce the
number of internal components and higher plates and masts that provide
additional plate area at a low cost. Microprocessor controls and energy
management systems have lowered operating costs.
Few, if any, hot-side ESPs (those used upstream from an air preheater
on a combustion source) are being specified for purchase. Recognition that
low sodium coals tend to build resistive ash layers on the collection plates,
thus reducing ESP efficiency, has almost eliminated sales of these units. Of
about 150 existing units, about 75 are candidates for conversion to cold-side
units over the next 10 years.
Specific industry application has little impact on either ESP design or
cost, with three exceptions: paper mills and sulfuric acid manufacturing
6-45
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plants, and coke by-product plants. Paper mill ESPs use drag conveyor
hoppers that add approximately 10 percent to the base flange-to-flange
equipment cost. For emissions control in sulfuric acid plants and coke by-
product ovens, wet ESPs are used. In sulfuric acid manufacture, wet ESPs
are used to collect acid mist. These precipitators usually are small, and
they use lead for all interior surfaces; hence, they normally cost $65 to
$95/ft2 of collecting area installed (mid-1987 dollars) and up to $120/ft2
in special situations. In addition, a wet circular ESP is used to control
emissions from a coke oven off-gas detarring operation. These precipitators
are made using high-alloy stainless steels and typically cost $90 to $120/ft2,
installed. Because of the small number of sales, small size of units sold, and
dependency on site-specific factors, more definitive costs are not available.
6.3.1.2 Retrofit Cost Factor
Retrofit installations increase the costs of an ESP because of the common
need to remove something to make way for the new ESP. Also, the ducting
usually is much more expensive. The ducting path is often constrained by
existing structures, additional supports are required, and the confined areas
make erection more labor intensive and lengthy. Costs are site-specific;
however, for estimating purposes, a retrofit multiplier of 1.3 to 1.5 applied to
the total capital investment can be used. The multiplier should be selected
within this range based on the relative difficulty of the installation.
A special case is conversion of hot-to-cold side ESPs for coal-fired boiler
applications. The magnitude of the conversion is very site-specific, but
most projects will contain the following elements:
• Relocating the air preheater and the ducting to it
• Resizing, the ESP inlet and outlet duct to the new air volume and
rerouting it
• Upgrading the ID (induced draft) fan size or motor to accommodate
the higher static pressure and horsepower requirements
• Adding or modifying foundations for fan and duct supports
• Assessing the required SCA and either increasing the collecting area
or installing an SO3 gas-conditioning system
6-46
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• Adding hopper heaters
• Upgrading the analog electrical controls to microprocessor-type con-
trols
• Increasing the number of collecting plate rappers and perhaps the
Location of rapping
In some installations, it may be cost-effective to gut the existing collector
totally, utilize only the existing casing and hoppers, and upgrade to modern
internals.
The cost of conversion is a multimillion dollar project typically running
at least 25 to 35 percent of the total capital investment of a new unit.
6.3.1.3 Auxiliary Equipment
The auxiliary equipment depicted in Figure 6.2 is discussed elsewhere in the
Manual. Because dust-removal equipment (e.g., screw conveyers), hoods,
precoolers, cyclones, fans, motors, and stacks are common to many pol-
lution control systems, they are (or will be) given extended treatment in
separate chapters.
6.3.1.4 Costs for Two-Stage Precipitators
Purchase costs for two-stage precipitators, which should be considered sep-
arately from large-scale, single-stage ESPs, are given in Figure 6.6.[7] To
be consistent with industry practice, costs are given as a function of flow
rate through the system. The lower cost curve is for a two-cell unit with-
out precooler, an installed cell washer, or a fan. The upper curve is for an
engineered, package system with the following components: inlet difFuser
plenum, prefilter, cooling coils with coating, coil plenums with access, water
flow controls, triple pass configuration, system exhaust fan with accessories,
outlet plenum, and in-place foam cleaning system with semiautomatic con-
trols and programmable controller. All equipment is fully assembled me-
chanically and electrically, and it is mounted on a steel structural skid.
6-47
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100
90
_ 80
i
| 70
•2 60
<• 50
I -
I"
30
20
10
-„+*
.Packaged Syttam
..Syttera Without
^ Praeootar. Imtaltod'
CaUWath^.orFan
0 8 10
Flow Raw (1.000 acfm)
12 14
Figure 6.6: Purchase Costs for Two-stage, Two-cell Precipitators[7]
6-48
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Table 6.8: Items That Increase ESP Costs
Item
Factor
Applied to
Rigid frame electrode with
restricted plate height
Type 304 stainless steel collector
plates and precipitator walls"
All stainless steel construction0
ESP with drag conveyor hoppers
(paper mill)
Retrofit installations
Wet ESP
Sulfuric acid mist
Sulfuric acid mist
(special installation)
Coke oven off gas
1.0 to 1.25 ESP base cost
1.3 to 1.5 "
2 to 3 "
1.1 "
1.3 to 1.5
ESP total capital
investment
/ new facility A
\ installation /
See 6.3.1.1.
See 6.3.1.1.
See 6.3.1.1.
'See table on page 6-45 for other materials' cost factors.
6.3.2 Total Purchased Cost
The total purchased cost of an ESP system is the sum of the costs of the
ESP, options, auxiliary equipment, instruments and controls, taxes, and
freight. The last three items generally are taken as percentages of the
estimated total cost of the first three items. Typical values, from Chapter
2 of the Manual, are 10 percent for instruments and controls, 3 percent for
taxes, and 5 percent for freight.
Costs of standard and other options can vary from 0 to more than 150
percent of bare ESP cost, depending on site and application requirements.
Other factors that can increase ESP costs are given in Table 6.8.
6-49
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6.3.3 Total Capital Investment (TCI)
Using the Chapter 2 methodology, TCI is estimated from a series of fac-
tors applied to the purchased equipment cost to obtain direct and indirect
costs for installation. The TCI is the sum of these three costs. The re-
quired factors are given in Table 6.9. Because ESPs may vary from small
units attached to existing buildings to large, separate structures, specific
factors for site preparation or for buildings are not given. However, costs
for buildings may be obtained from such references as Means Square Foot
Costa 1987 [10]. Land, working capital, and off-site facilities are excluded
from the table because they are not normally required. For very large in-
stallations, however, they may be needed and could be estimated on an
as-needed basis.
Note that the factors given in Table 6.9 are for average installation con-
ditions, e.g., no unusual problems with site earthwork, access, shipping, or
interfering structures. Considerable variation may be seen with other-than-
average installation circumstances. For two-stage precipitators purchased
as packaged systems, several of the costs in Table 6.9 would be greatly re-
duced or eliminated. These include instruments and controls, foundations
and supports, erection and handling, painting, and model studies. An in-
stallation factor of 0.20 B to 0.25 B would be more nearly appropriate for
the two-stage ESPs.
6.4 Estimating Total Annual Costs
6.4.1 Direct Annual Costs
Direct annual costs include operating and supervisory labor, operating ma-
terials, replacement rappers and electrodes, maintenance (labor and ma-
terials), utilities, dust disposal, and wastewater treatment for wet ESPs.
Most of these costs are discussed individually below. They vary consid-
erably with location and time and, for this reason, should be obtained to
suit the specific ESP system being costed. For example, current labor rates
may be found in such publications as the Monthly Labor Review, published
by the U.S. Department of Labor, Bureau of Labor Statistics.
6-50
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Table 6.9: Capital Cost Factors for ESPs"
Cost Item Factor
Direct Costs
Purchased equipment costs
ESP + auxiliary equipment As estimated, A
Instrumentation 0.10 A
Sales taxes 0.03 A
Freight 0.05 A
Purchased equipment cost, PEC B = 1.18 A
Direct installation costs
Foundations & supports 0.04 B
Handling & erection 0.50 B
Electrical 0.08 B
Piping 0.01 B
Insulation for ductwork* 0.02 B
Painting 0.02 B
Direct installation costs 0.67 B
Site preparation As required, SP
Buildings As required, Bldg.
Total Direct Costs, DC 1.67 B + SP + Bldg.
Indirect Costs (installation)
Engineering 0.20 B
Construction and field expenses 0.20 B
Contractor fees 0.10 B
Start-up 0.01 B
Performance test 0.01 B
Model study 0.02 B
Contingencies 0.03 B
Total Indirect Costs, 1C '0.57 B
Total Capital Investment = DC + 1C 2.24 B + SP + Bldg.
•Reference [8]
*If ductwork dimensions have been established, cost may be estimated based on $10 to $12/ft3
(fourth quarter 1986) of surface for field application. Fan housings and stacks may also be
insulated.[0j
'For two-stage precipitators, total installation direct, costs are more nearly 0.20 to 0.25B + SP +
Bldg..
6-51
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6.4.1.1 Operating and Supervisory Labor
Proper operation of the ESP is necessary both to meet applicable particu-
late emission regulations and to ensure minimum costs. An ESP is an ex-
pensive piece of equipment. Even well-designed equipment will deteriorate
rapidly if improperly maintained and will have to be replaced long before it
should be necessary. Not only can proper operation and maintenance save
the operator money, such an operation and maintenance program can also
contribute to good relations with the governing pollution control agency by
showing good faith in efforts to comply with air regulations.
Although each plant has its own methods for conducting an operation
and maintenance program, experience has shown that plants that assign one
individual the responsibility of coordinating all the pieces of the program
operate better than those where different departments look after only a cer-
tain portion of the program. The separate departments have little knowl-
edge of how their portion impacts the overall program. In other words, a
plant needs one individual to coordinate the operation, maintenance, and
troubleshooting components of its ESP program if it expects to have a
relatively trouble-free operation. The coordinator typically is an engineer
who reports to plant management and interfaces with the maintenance and
plant process supervisors, the laboratory, and the purchasing department.
For companies with more than one plant, he would be responsible for all
ESPs. The portion of his total time that this individual spends on the ESP
then becomes an operating expense for the ESP. This can be expressed as:
AC = X(LCC) (6.25)
where
AC = annual coordination cost ($/yr)
X = fraction of total time spent on ESP
LCC = individual annual labor cost for ESP coordinator ($/yr)
In addition to coordination costs, typical operating labor requirements are
1/2 to 2 hours per shift for a wide range of ESP sizes.[8] Small or well-
performing units may require less time, and very large or troublesome units
may require more time. Supervisory labor is taken as 15 percent of oper-
ating labor.
6-52
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6.4.1.2 Operating Materials
Operating materials are generally not required for ESPs. An exception is
the use of gas-preconditioning agents for dust resistivity control.,
6.4.1.3 Maintenance
The reader should obtain Publication No. EPA/625/1-85/017, Operating
and Maintenance Manual for ESPa,[ll] for suggested maintenance prac-
tices. Routine ESP maintenance labor costs can be estimated using data
provided by manufacturers. If such data is unavailable, the following pro-
cedure can be used. Based on data for a 100,000 ft2 collector, maintenance
labor is estimated to require 15 h/wk, 44 wk/yr. At a direct labor cost
of $12.50/h (mid-1987 costs), an estimated annual maintenance labor cost
of $8,250 or $0.0825/ft2 of collector area is established. This relationship
can be assumed to be linear above a 50,000 ft2 collector-size and constant
at $4,125 below this size. To the maintenance labor cost must be added
the cost of maintenance materials. Based on an analysis of vendor infor-
mation, annual maintenance materials are estimated as 1 percent of the
flange-to-flange precipitator purchase cost:
MC = QM(FCC) + labor cost (6.26)
where
MC = annual maintenance cost ($/yr)
FCC = ESP flange-to-flange purchase cost ($)
, , . / $4,125 if A < 50,000 ft2
labor cost = < , .. . ..,
0.08254 if A > 50,000 ft2 ($)
where A = ESP plate area (ft2)
6.4.1.4 Electricity
Power is required to operate system fans, transformer-rectifier(TR) sets,
and cleaning equipment. Fan power for primary gas movement can be
calculated from Equation 2.7 of the Manual, After substituting into this
equation a combined fan-motor efficiency of 0.65 and a specific gravity of
1.0, we obtain:
FP = 0.000181C?(AP)(0') (6.27)
6-53
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where
FP = fan power requirement (kWh/yr)
Q — system flow rate (acfm)
AP = system pressure drop (in. HjO)
tf = annual operating time (h/yr)
Pump power for wet ESPs can be calculated from [8]:
PP = (0.746 Qi Z Sg 0')/(3,960?;) (6.28)
where
PP = pump power requirement (kWh/yr)
Qi = water flow rate (gal/min)
Z = fluid head (ft)
Sg = specific gravity of water being pumped compared to
water at 70 °F and 29.92 in. Hg
6' = annual operating time (h/yr)
rj = pump-motor efficiency (fractional)
Energy for TR sets and motor-driven or electromagnetic rapper systems
is the sum of the energy consumption for operating both" items. Manufac-
turers' averaged data indicate that the following relationship can be used:
OP = 1.94 x IQ-3A9' (6.29)
where
OP = annual ESP operating power (kWh/yr)
A = ESP plate area (ft2)
ff = annual operating time (h/yr)
For installations requiring hopper heaters, hopper heater power can be
similarly estimated:
HH = 2(HN)6' (6.30)
where
HH = annual hopper heater power consumption (kWh/yr)
HN = number of hoppers
0' = annual operating time (h/yr)
For two-stage precipitators, power consumption ranges from 25 to 100
W/kacfm, with 40 W/kacfm being typical.
6-54
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6.4.1.5 Fuel
If the ESP or associated ductwork is heated to prevent condensation, fuel
costs should be calculated as required. These costs can be significant, but
they may be difficult to predict. For methods of calculating heat transfer
requirements, see Perry [12].
6.4.1.6 Water
Cooling process gases for preconditioning can be done by dilution with air,
evaporation with water, or heat exchange with normal equipment. Spray
cooling requires consumption of plant water (heat exchangers may also
require water), although costs are not usually significant. Section 4.4 of
the Manual provides information on estimating cooling water costs. Water
consumption in wet ESPs is estimated at 5 gal/min kacfm [13] for large
single-stage units and 16 gal/min-kacfm for two-stage precipitators [14].
6.4.1.7 Compressed Air
*
ESPs may use compressed air at pressures of about 60 to 100 psig for
operating rappers. Equivalent power cost is included in Equation 6.29 for
operating power.
6.4.1.8 Dust Disposal
If collected dust cannot be recycled or sold, it must be landfilled or disposed
of in some other manner. Costs may typically run $20/ton or $30/ton for
nonhazardous wastes exclusive of transportation (see Section 2.4 of the
Manual). Landfilling of hazardous wastes may cost 10 times as much. The
disposal costs are highly site-specific and depend on transportation distance
to the landfill, handling rates, and disposal unloading (tipping) fees. If these
factors are known, they lead to the relationship:
DD = 4.29 x W~9G ff Q(T + (TM)D] (6.31)
6-55
-------
where
DD = annual dust disposal cost ($/yr)
G = ESP inlet grain loading or dust concentration (gr/ft3)
9' = annual operating time (h/yr)
Q = gas flow rate through ESP (acfm)
T = tipping fee ($/ton)
TM = mileage rate ($/ton-mile)
D = dust hauling distance (miles)
6.4.1.9 Wastewater Treatment
As indicated above, the water usage for wet ESPs is about 5 gal/min kacfm
[13]. Treatment cost of the resulting wastewater may vary from about
$1.30 to $2.15/1,000 gal [15] depending on the complexity of the treatment
system. More precise costs can be obtained from Gumerman et al [16].
6.4.1.10 Conditioning Costs
Adaptation of information on utility boilers [17] suggests that S03 condi-
tioning for a large ESP (2.6 x 10a acfm) costs from about $1.60/106 ft3 of
gas processed for a sulfur burner providing 5 ppm of S03 to about $2.30/10
ft3 (in first-quarter 1987 dollars) for a liquid S02 system providing 20 ppm
ofS03.
6.4.2 Indirect Annual Costs
Capital recovery, property taxes, insurance, administrative costs ("G&A"),
and overhead are examples of indirect annual costs. The capital recovery
cost is based on the equipment lifetime and the annual interest rate em-
ployed. (See Chapter 2 for a thorough discussion of the capital recovery
cost and the variables that determine it.) For ESPs, the system lifetime
varies from 5 to 40 years, with 20 years being typical. Therefore, as Chapter
2 of the Manual suggests, when figuring the system capital recovery cost,
one should base it on the total capital investment. In other words:
CRCs = TCI x CRFs (6.32)
6-56
-------
where
CRCa = capital recovery cost for ESP system ($/yr)
TCI = total capital investment ($)
CRFs — capital recovery factor for ESP system (defined in
Chapter 2)
For example, for a 20-year system life and a 10 percent annual interest rate,
the CRFs would be 0.1175.
The suggested factor to use for property taxes, insurance, and adminis-
trative charges is 4 percent of the TCI. Overhead is calculated as 60 percent
of the sum of operating, supervisory, coordination, and maintenance labor,
as well as maintenance materials.
6.4.3 Recovery Credits
For processes that can reuse the dust collected in the ESP or that can sell
the dust in a local market, such as fly ash sold as an extender for paving
mixes, a credit should be taken. As used below, this credit (RC) appears
as a negative cost.
6.4.4 Total Annual Cost
Total annual cost for owning and operating an ESP system is the sum of
the components listed in Subsections 6.4.1 through 6.4.3, i.e.:
TAG = DC + IC-RC (6.33)
where
TAC = total annual cost ($)
DC = direct annual cost ($)
1C = indirect annual cost ($)
RC = recovery credits (annual) ($)
6-57
-------
6.4.5 Example Problem
Assume an ESP is required for controlling fly ash emissions from a coal-
fired boiler burning bituminous coal. The flue gas stream is 50 kacfm at
325°F and has an inlet ash loading of 4 gr/ft3. Analysis of the ash shows
an MMD of 7 /*m and a resistivity of less than 2 x 1011 ohm-cm. Assume
that the ESP operates for 8,640 h/yr (360 d) and that an efficiency of 99.9
percent is required.
6.4.5.1 Design SCA
The SCA can be calculated from Equation 6.23. Assuming that a flat plate
ESP design is chosen, the fly ash migration velocity is 16.0 cm/s (see Table
6.4). Then:
SCA = -ln(l - 0.999)/16.0 = 0.432 s/cm = 43.2 s/m
Converting to English units (see page 6-33, Step 15, in the procedure):
ESCA = 5.080 x 43.2 = 219 ft2/kacfm
Total collector plate area is then:
219 ft'/kacfm x 50 kacfm = 10,950 ft2
To obtain a more rigorous answer, we can follow the steps of the proce-
dure given in Subsection 6.2.1:
Step 1 - Design efficiency is required as 99.9.
Step 2 - Design penetration:
1 - (99.9/100) = 0.001
6-58
-------
Step 3 - Operating temperature in Kelvin :
(325°F - 32°.F) x 5/9 + 273°C = 436A"
Step 4 - Because dust resistivity is less than 2 x 1011 ohm-cm (see page
6-30, Step 4), no severe back corona is expected and back corona = 0.
Step 5 - The MMD of the fly ash is given as 7
Step 6 - Values for sneakage and rapping reentrainment (from the table
presented in Step 6, page 6-30) are:
8* = 0.10
RR = 0.124 (assuming gas velocity <1.5 m/s)
Step 7 - The most penetrating particle size, from Step 7 of the procedure
on page 6-31, is:
MMDp = 2/im
The rapping puff size is:
MMDP = 5/zm
Step 8 - From the procedure (Subsection 6.2.1):
eO = 8.845 xlO'12
r, = 1.72 x 10-5(436/273)°-71 = 2.40 x lO"5
Ew = 6.3 x 105(273/436)1-6B = 2.91 x 106 V/m
E00ff = EM x 5/6.3 = 2.31 x 105
LF = 8* + RR(1 - $N) = 0.1 + 0.124(1 - 0.1) = 0.212
6-59
-------
Step 9 - Choose the number of sections for LFn < p, p = 0.001. Try four
sections:
LF" = 0.2124 = 0.002
This value is larger than p. Try five sections:
IF" = 0.2128 = 0.000428
This value is smaller than p and is acceptable.
Step 10 - Average section penetration is:
p, = p1/" = 0.0011'5 = 0.251
Step 11 - Section collection penetration is:
Pc = (p. - LF)/(l - LF) = (0.251 - 0.212)/(1 - 0.212) = 0.0495
Step 12 - Particle size change factors are:
D = p, = SN + Pe(l - SN) + RR(l - SN)(l - PC)
= 0.10 + 0.0495(1 - 0.1) + 0.124(1 - 0.1)(1 - 0.0495)
= 0.251
MMDrp = RR(l -Sff)(l-pc)MMDr/D
= 0.124(1 - 0.1)(1 - 0.0495)(5)/0.251
= 2.11
Step 13 - Particle sizes for each section are:
6-60
-------
Section MMD (/mi)
= MMDi = 7
2 MMD2 = {MMDjX SN + [(I - pe) x MMDP + Pex
MMDi] xpj/D + MMDrp
= {7 x 0.1 + [(1 - 0.0495) x 2 + 0.0495 x 7]
x 0.0495}/0.251 + 2.11
= 5.34
3 MMD3 = {5.34x0.1 + [(l-0.0495)x2+0.0495x5.34]x
0.0495 }/0.251 + 2.11
= 4.67
4 MMD4 = {4.67x0.1 + [(l-0.0495)x2+0.0495x4.67]x
0.0495 }/0.251 + 2.11
= 4.39
5 MMD5 = {4.39x0.1 + [(l-0.0495)x2+0.0495x4.39]x
0.0495 }/0.251 + 2.11
= 4.28
Step 14 - SCAs for each section are:
Section SCA (s/m)
1 SCA: = -(77/eO)x(l-5*)xln(p^/(£^xMMD1xlO-8)
= -(2.40 x 10~8/8.845 x lQ-n)(l - 0.1) x
ln(0.0495)/[(2.31 x 10B)2(7 x lO'8)]
= 19.65
2 SCA2 = SCAiX MMD!/MMD2
= 19.65 (7/5.34)
= 25.76
3 SCA3 = 25.76(5.34/4.67)
= 29.46
4 SCA4 = 29.46(4.67/4.39)
= 31.34
5 SCA5 = 31.34(4.39/4.28)
= 32.15
Step 15 - Calculate the total SCA.
6-61
-------
Total SCA = 19.65 + 25.76 + 29.46 + 31.34 + 32.15 = 138.36 s/m
English SCA = 5.080 x 138.36 = 702.87 ft2/kacfm
Note that the more rigorous procedure calls for an SCA that is consider-
ably higher than the value found by using Equation 6.23. This discrepancy
is caused by the considerably smaller particle size used in the example prob-
lem than is assumed for Table 6.4. In this case, the shorter method would
lead to an unacceptably low cost estimate.
Total collector plate area is:
702.87 ft'/kacfm x 50 kacfm = 35,144 ft2
6.4.5.2 ESP Cost
From Figure 6.5, the basic flange-to-flange cost of the rigid electrode ESP is
$438,060 (mid-1987 dollars). Assuming all standard options are purchased,
the ESP cost rises to $635,189 (mid-1987 dollars).
6.4.5.3 Costs of Auxiliaries
Assume the following auxiliary costs have been estimated from data in other
parts of the Manual:
Ductwork $16,000
Fan 16,000
Motor 7,500
Starter 4,000
Dampers 7,200
Pneumatic conveyor 4,000
Stack 8,000
Total $62,700
6-62
-------
6.4.5.4 Total Capital Investment
Direct costs for the ESP system, based on the factors in Table 6.9, are given
in Table 6.10. (Again, we assume site preparation and building costs to be
negligible.) TCI is $1,840,000 (rounded, mid-1987 dollars).
6.4.5.5 Annual Costs-Pressure Drop
Table 6.11 gives the direct and indirect annual costs, as calculated from
the factors given in Section 6.4. Pressure drop (for energy costs) can be
taken from Table 6.6 in Subsection 6.2.2. Using the higher values from the
table, pressure drop for the inlet diffuser plate, inlet and outlet transitions,
baffles, and plates is:
AP = 0.09 + 0.14 + 0.015 + 0.123 + 0.008 = 0.38 in. H2O
Assume the ductwork contributes an additional 4.1 in. H20. The total pres-
sure drop is, therefore, 4.48 in. H20. As is typical, the ductwork pressure
drop overwhelms the ESP pressure drop.
6.4.5.6 Total Annual Cost
The total annual cost, calculated in Table 6.11, is $553,000 (rounded). Had
the particle sizes being captured been larger, the ESP cost would, have been
considerably less. Also, for a much larger gas flow rate, the $/acfm treated
cost would have been more favorable. Reviewing components of the TAG,
dust disposal is the largest single item. Care should be taken in determining
this cost and the unit disposal cost ($/ton). Finding a market for the dust,
for example, as an extender in asphalt or a dressing for fields, even at give
away prices, would reduce TAG dramatically.
6-63
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Table 6.10: Capital Costs for ESP System
Example Problem
Cost Item Cost
Direct Costs
Purchased equipment costs
Adsorber vessels and carbon $635,189
Auxiliary equipment 62,700
Sum = A $697,889
Instrumentation, 0.1A 69,789
Sales taxes, 0.03A 20,937
Freight, 0.05A 34,894
Purchased equipment cost, B $823,509
Direct installation costs
Foundation and supports, 0.04B 32,940
Handling & erection, 0.50B 411,755
Electrical, 0.08B 65,881
Piping, 0.01B 8,235
Insulation for ductwork, 0.02B 16,470
Painting, 0.02B 1MI°_
Direct installation cost $551,751
Site preparation
Facilities and buildings
Total Direct Cost $1,375,260
Indirect Costs (installation)
Engineering, 0.20B 164,702
Construction and field expenses, 0.20B 164,702
Contractor fees, 0.10B 82,351
Start-up, 0.01B 8,235
Performance test, 0.01B 8,235
Model study, 0.02B 16,470
Contingencies, 0.03B 24,705
Total Indirect Cost $469,400
Total Capital Investment (rounded) $1,840,000
6-64
-------
Table 6.11: Annual Costs for ESP System
Example Problem
Cost Item
Calculations
Cost
Direct Annual Costs, DC
Operating labor
Operator
Supervisor
Coordinator
Operating materials
Maintenance
Labor
Material
Utilities
Electricity-fan
Electricity-operating
Waste disposal
360 days $12
— yr XT
15% of operator = .15 x 12,960
1/3 of operator = 1/3 X 12,960
$ 4,125 for collector area < 50,000 ft2
1% of Purchased equipment cost = 0.01 x 823,509
0.000181 x 50,000 acfm x 4.48 in. H2O x 8'6yr° x
$0.06
kWh
1.94 x 10-3 x 35,144.fts x 8,640 h x $0.06/kWh
at $20/ton tipping fee at 2 miles and $0.50/ton-
mile for essentially 100% collection efficiency: 4.29 x
10
-«
8'6yir0 h x 50,000 acfm x (20 + 0.50 x
$12,960
1,944
4,320
4,125
8,235
21,018
35,344
155,676
Total DC
$243,622
Indirect Annual Costs, 1C
Overhead 60% of sum of operating, supv., coord., & maint. 18,950
labor & maint. materials = 0.6(12,960 + 1,944 +
4,320 + 4,125 + 8,235)
Administrative charges 2% of Total Capital Investment = 0.02($1,844,660) 36,893
Property tax 1% of Total Capital Investment = 0.01($1,844,660) 18,447
Insurance 1% of Total Capital Investment = 0.01($1,844,660) 18,447
Capital recovery" 0.1175 (1,844,660) 216,748
Total 1C *309,485
Total Annual Cost (rounded)
$553,000
0 The capital recovery cost factor, CRF, is a function of the fabric filter or equipment life
and the opportunity cost of the capital (i.e., interest rate). For example, for a 20 year
equipment life and a 10% interest rate, CRF = 0.1175.
6-65
-------
6.5 Acknowledgments
We gratefully acknowledge C. G. Noll, United McGill Corp. (Columbus,
OH), for extensive review and the following companies for contributing data
to this chapter:
• Research- Cottrell
• Joy Industrial Equipment Co., Western Precipitation Division (Los
Angeles, CA)
• Environmental Elements Corp. (Baltimore, MD)
6-66
-------
-------
Appendix 6A
Effects of Material Thickness
and Type on ESP Costs
The impact of material thickness and composition of collecting plates and
the ESP casing can be estimated using the following equations and Figure
6.5:
Plates:
[(^ x FS) - 0.90] M + SP
Casing:
(6.34)
_ [(2£ x F5) - 0.58] M + SP
j= _
6-67
-------
where
I = incremental increase of flange-to-flange selling price
($/ft2)
Wt = weight of steel (lb/ft2)
FS = fabricated steel selling price ($/lb) (normally assume
approximately 2 times material cost)
M = manufacturer's markup factor of fabricated cost (direct
labor, wages, and material cost before general and ad-
ministrative expense and profit) to selling price (nor-
mally 2 to 3)
SP = flange-to-flange selling price from Figure 6.5 ($/ft2)
Most vendors can produce ESPs with collecting plate material thick-
nesses from 16 to 20 gauge and casing material thicknesses from 1/8 through
1/4 in. without affecting the 2 times material cost = fabricated cost rela-
tionship. Thus, the impact of increasing the collecting plates from 18 to 16
gauge and the casing from 3/16 to 1/4 in. plate on a 72,000 ft2 collector
having a selling price of $10/ft2 and assuming a markup factor of 2 is as
follows:
Plates:
[( Y x 0.90) - 0.90] 2 + 10
= _
= 1.045 = 4.5 percent increase
0.76) - 0.58] 2 + 10
/ = —
= 1.039 = 3.9 percent increase
Equations 6.34 and 6.35 were developed using the following assumptions:
_ Material selling price increase + Standard ESP selling price
1 = Standard ESP selling price
6-68
-------
Figure 6.7: ESP Dimensions
Because Figure 6.5 identifies the standard ESP selling price/ft2 of collecting
area, the material selling price increase = (New material cost - Standard
material cost)M. Then it follows that:
Ib steel
Material selling price = —r-» = : x Fabricated cost in $/lb X M
fir collecting area
The ESP dimensions given in Figure 6.7 include:
• Casing area = 30 ft x 30 ft x 8 = 7,200 ft2 (assume 4 walls, 1 top, 2
hopper sides, 2 triangular hopper ends =s 8 equivalent sides)
6-69
-------
• Collecting plate area =
30 ft x 30 ft x 2 sides/plate x plates
3
= 54'000 ft* = 72,000 ft2 for a = 0.75 ft
where a = plate spacing (ft)
Thus, there are:
• 7.50/a ft2 of collecting area per 1 ft2 of casing and
• 2 ft2 of collecting area per 1 ft2 of collecting plate
Material cost per ft2 collecting area is:
Ib steel/ft2 , ...
Plates s= —i x $/lb
Ib steel/ft2 « .
7.50/s X $/lb
For the standard ESPs described by Figure 6.5, 18 gauge collecting plates
and 3/16 in. plate casing were specified. Assuming:
Material cost for 18 gauge mild steel = $0.45/lb
Material cost for 3/16 in. plate mild steel = $0,38/lb
Material cost to fabricated cost factor = 2
These costs yield fabricated material costs of:
Plates:
6-70
-------
2 lb/ft2 x $0.45/lb x 2 = $0.90/ft2 of collecting area
Casing:
•_ >K^fe 'H / *i 4
x $0.38 x 2 = $0.78 a/ft2 of collecting area
_.
I •O\J f 3
At a typical 9 in. plate spacing the casing cost would be $0.58/ft2 of col-
lecting area.
Selling / Cost of _ Cost of \ M Original overaU
price = V new material old material / selling price
impact Original overall selling price
which gives us equations 6.34 and 6.35. Note that the value 0.58 will change
significantly if a plate spacing other than 9 in. is chosen.
Thus, for a less than 5 percent increase in flange-to-flange cost, all the
precipitator exposed wall sections can be increased by more than 25 per-
cent to provide increased life under corrosive conditions. Section thickness
increases that are greater than those just discussed would probably result
in significant cost increases because of both increased material costs and
necessary engineering design changes.
The impact of changing from mild steel to 304 stainless steel assuming
material costs of $1.63/lb for 18 gauge collecting plates, $1.38/lb for the
3/16 in. casing, and a markup factor of 3 is as follows:
Plates:
[(I x 1.63) - 0.9] 3 + 10
1 = 10
= 21.9 percent increase
[(2jK x 1.38) - 0.58] 3 + 10
10
6-71
-------
= 14.3 percent increase
To these material costs must be added extra fabrication labor and pro-
curement costs that will increase the ESP flange-to-flange cost by a factor
of 2 to 3. Note that a totally stainless steel collector would be much more
expensive because the discharge electrodes, rappers, hangers, etc., would
be also converted to stainless. The preceding equations can be used for
other grades of stainless steel or other materials of construction by insert-
ing material costs obtained from local vendors on a $/lb basis.
6-72
-------
References
[1] White, H. J., Industrial Electrostatic Precipitation, Addison-Wesley,
Reading, MA, 1963.
[2] Lawless, P. A., and L. E. Sparks, "A Review of Mathematical Models
for ESPs and Comparison of Their Successes," Proceedings of Second
International Conference on Electrostatic Precipitation, S. Masuda,
ed., Kyoto, 1984, pp. 513-522.
[3] Bump, R. L. (Research Cottrell, Inc.), "Evolution and Design of Elec-
trostatic Precipitator Discharge Electrodes," paper presented at the
APCA Annual Meeting, New Orleans, LA, June 1982.
[4] Correspondence: Richard Selznick (Baron Blakeslee, Inc., Westfield,
NJ) to William M. Vatavuk, April 23, 1986.
[5] Correspondence: James Jessup (M&W Industries, Inc., Rural Hall,
NC) to William M. Vatavuk, May 16, 1986.
[6] Matley, Jay (ed.), Modern Cost Engineering, McGraw-Hill, New York,
1984, p. 142.
[7] Personal communication: Robert Shipe, Jr. (American Air Filter Co.,
Louisville, KY), and S. A. Sauerland (United Air Specialists, Inc.,
Cincinnati, OH), to Roger Ellefson (JACA Corp., Fort Washington,
PA), June 1987.
[8] Vatavuk, W. M., and R. B. Neveril, "Estimating Costs of Air Pollution
Control Systems, Part II: Factors for Estimating Capital and Operat-
ing Costs," Chemical Engineering, November 3, 1980, pp. 157-162.
[9] Telecon: Gary Greiner (ETS, Inc., Roanoke, VA) to James H. Turner,
October 1986.
6-73
-------
[10] R. S. Means Company, Inc., Means Square Foot Coats 1987, Kingston,
MA.
[11] PEDCo Environmental, Inc., Operating and Maintenance Manual for
ESPs, Publication No. EPA/625/1-85/017, Office of Research and De-
velopment, Air and Energy Engineering Research Lab, Research Tri-
angle Park, NC, September 1985.
[12] Perry, R. H., et aL, Perry's Chemical Engineers' Handbook (Sixth Edi-
tion), McGraw-Hill, New York, 1984.
[13] Bakke, E., "Wet Electrostatic Precipitators for Control of Sub-micron
Particles," Proceedings of the Symposium on Electrostatic Precipitators
for the Control of Fine Particles, Pensacola, FL, September 30 to
October 2, 1974, Publication No. EPA-650/2-75-016, 1975.
[14] Beltran Associates, Inc., "Poly-Stage Precipitator for Stack and Duct
Emissions," November 1978.
[15] Vatavuk, W. M., and R. B. Neveril, "Estimating Costs of Air-Pollution
Control Systems, Part XVII: Particle Emissions Control," Chemical
Engineering (adapted), April 2, 1984, pp. 97-99.
[16] Gumerman, R. C., B. E. Burns, and S. P. Hansen, Estimation of Small
System Water Treatment Costs, Publication No. EPA/600/2-84/184a,
NTIS No. PB85-161644, 1984.
[17] Gooch, J. P., A Manual on the Use of Flue Gas Conditioning for ESP
Performance Enhancement, Electric Power Research Institute Report
No. CS-4145, 1985.
6-74
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Chapter 7
FLARES
Diana K. Stone
Susan K. Lynch
Richard F. Pandullo
Radian Corporation
Research Triangle Park, NC 27709
Leslie B. Evans, Chemicals and Petroleum Branch
William M. Vatavuk, Standards Development Branch
Office of Air Quality Planning and Standards
U.S. Environmental Protection Agency
Research Triangle Park. NC 27711
April 1991
Contents
7.1 Introduction 7-4
7.1.1 Flare Types 7-4
7-1
-------
7.1.1.1 Steam-Assisted Flares .............. 7.5
7.1.1.2 Air-Assisted Flares ............... 7.5
7.1.1.3 Non-Assisted Flares ............... 7.6
7.1.1.4 Pressure- Assisted Flares ............ 7-6
7.1.1.5 Enclosed Ground Flares ............ 7.5
7.1.2 Applicability ........................ 7_7
7.1.3 Performance ....................... 7,g
•\»
7.1.3.1 Factors Affecting Efficiency .......... 7.3
7.1.3.2 Flare Specifications ............... 7.9
7.2 Process Description .................. 7-10
7.2.1 Gas Transport Piping ................... 7. 10
7.2.2 Knock-out Drum ............... 7-10
7.2.3 Liquid Seal ..................... 7,0
7.2.4 Flare Stack ......................... 7_13
7.2.5 Gas Seal .......................... 7_17
7.2.6 Burner Tip .................. 7,17
7.2.7 Pilot Burners
7.2.8 Steam Jets
7 ,«-
7.2.9 Controls ....................... 7,ig
7.3 Design Procedures .................... 7 19
7.3.1 AmrW — Fuel Requirement ............... 7_19
7.3.2 Flare Tip Diameter
7.20
7-2
-------
7.3.3 Flare Height 7.22
7.3.4 Purge Gas Requirement 7_23
7.3.5 Pilot Gas Requirement 7.24
7.3.6 Steam Requirement -
-------
7.1 Introduction
Flaring is a volatile organic compound (VOC) combustion control process in
whichJji^VOCs are pipe? toa remote, usually elevated, location andburned
in an open flame in the open air using a specially designed burner tip, aux-
ilija.rjrjuel^an.d steam or air to promotejnixing for nearlycomplete (> 98%)
_VOC_destructipn. Completeness of combustion in a flare is governed by flame
temperature, residence time in the combustion zone, turbulent mixing of the
components to complete the oxidation reaction, and available oxygen for free
radical formation. Combustion is complete if all VOCs are converted to car-
bon dioxide and water. Incomplete combustion results in some of the VOC
being unaltered or converted to other organic compounds such as aldehydes
or acids.
The flaring process can produce some undesirable by-products includ-
ing noise, smoke, heat radiation, light, SOX, NOX, CO, and an additional
source of ignition where not desired. However, by proper design these can be
minimized.
7.1.1 Flare Types
Flares are generally categorized-in two ways: (1) by the height of the flare tip
(i.e., ground pr elevated), and (2) by the method of enhancing mixing at the
flare tip (t.e., steam-assisted, air-assisted, pressure-assisted, or non-assisted).
Elevating the flare can prevent potentially dangerous conditions at ground
level where the open flame (i.e., an ignition source) is located near a process
unit. Further, the products of combustion can be dispersed above working
areas to reduce the effects of noise, heat, smoke, and objectionable odors.
In most flares, combustion occurs by means of a diffusion flame. A diffu-
sion flame is one in which air diffuses across the boundary of the fuel/combus-
tion product stream toward the center of the fuel flow, forming the envelope
of a combustible gas mixture around a core of fuel gas. This mixture, on ig-
nition, establishes a stable flame zone axound the gas core above the burner
tip. This inner gas core is heated by diffusion of hot combustion products
from the flame zone.
Cracking can occur with the formation of small hot particles of carbon
7-4
-------
that give the flame its characteristic luminosity. If there is an oxygen de-
ficiency and if the carbon particles are cooled to below their ignition tem-
perature, smoking occurs. In large diffusion flames, combustion product
vortices can form around burning portions of the gas and shut off the supply
of oxygen. This localized instability causes flame flickering, which can be
accompanied by soot formation.
As in all combustion processes, an adequate air supply and good mixing
are required to complete combustion and minimize smoke. The various flare
designs differ primarily in their accomplishment of mixing.
7.1.1.1 Steam-Assisted Flares
Steam-assisted flares are single burner tips; elevated above ground level for
safety reasons, that burn the vented gas in essentially a diffusion flame. They
reportedly account for the majority of the flares installed and are the pre-
dominant flare type found in refineries and chemical plants.[1, 2]
To ensure an adequate air supply and good mixing, this type of flare
system injects steam into the combustion zone to promote turbulence for
mixing and to induce air into the flame. Steam-assisted flares are the focus of
the chapter and will be discussed in greater detail in Sections 7.2 through 7.4.
7.1.1.2 Air-Assisted Flares
Some flares use forced air to provide the combustion air and the mixing
required for smokeless operation. These flares are built with a spicier-shaped
burner (with many small gas orifices) located inside but near the top of a steel
cylinder two feet or more in diameter. Combustion air is provided by a fan
in the bottom of the cylinder. The amount of combustion air can be varied
by varying the fan speed. The principal advantage of the air-assisted flares
is that they can be used where steam is not available. Although air assist is
not usually used on large flares (because it is generally not economical when
the gas volume is large(3j) the number of large air-assisted flares being built
is increasing.[4]
7-5
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7.1.1.3 Non-Assisted Flares
The non-assisted flare is just a flare tip without any auxiliary provision for
enhancing the mixing of air into its flame. Its use is limited essentially to
gas streams that have a low heat content and a low carbon/hydrogen ratio
that burn readily without producing smoke.[5] These streams require less air
for complete combustion, have lower combustion temperatures that minimize
cracking reactions, and are more resistant to cracking.
7.1.1.4 Pressure-Assisted Flares
Pressure-assisted flares use the vent stream pressure to promote mixing at
the burner tip. Several vendors now market proprietary, high pressure drop
burner tip designs. If sufficient vent stream pressure is available, these flares
can be applied to streams previously requiring steam or air assist for smoke-
less operation. Pressure-assisted flares generally (but not necessarily) have
the burner arrangement at ground level, and consequently, must be located
in a remote area of the plant where there is plenty of space available. They
have multiple burner heads that are staged to operate based on the quantity
of gas being released. The size, design, number, and group arrangement of
the burner heads depend on the vent gas characteristics.
7.1.1.5 Enclosed Ground Flares
An enclosed flare's burner heads are inside a. shell that is internally insu-
lated. This shell reduces noise, luminosity, and heat radiation and provides
wind protection. A high nozzle pressure drop is usually adequate to provide
the mixing necessary for smokeless operation and air or steam assist is not
required. In this context, enclosed flares can be considered a special class
of pressure-assisted or non-assisted flares. The height must be adequate for
creating enough draft to supply sufficient air for smokeless combustion and
for dispersion of the thermal plume. These flares are always at ground level.
Enclosed flares generally have less capacity than open flares and are used
to combust continuous, constant flow vent streams, although reliable and ef-
ficient operation can be attained over a wide range of design capacity. Stable
combustion can be obtained with lower Btu content vent gases than is possi-
7-6
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ble with open flare designs (50 to 60 Btu/scf has been reported)[2], probably
due to their isolation from wind effects. Enclosed flares are typically found
at landfills.
7.1.2 Applicability
Flares can be used to control almost any VOC stream, and can handle fluc-
tuations in VOC concentration, flow rate, heating value, and inerts content.
Flaring is appropriate for continuous, batch, and variable flow vent stream
applications. The majority of chemical plants and refineries have existing
flare systems designed to relieve emergency process upsets that require re-
lease of large volumes of gas. These large diameter flares, designed to handle
emergency releases, can also be used to control vent streams from various
process operations. Consideration of venlt stream flow rate and available
pressure must be given for retrofit applications. Normally, emergency relief
flare systems are operated at a small percentage of capacity and at negligi-
ble pressure. To consider the effect of controlling an additional vent stream,
the maximum gas velocity, system pressure, and ground level heat radiation
during an emergency release must be evaluated. Further, if the vent stream
pressure is not sufficient to overcome the flare system pressure, then the eco-
nomics of a gas mover system must be evaluated. If adding the vent stream
causes the maximum velocity limits or ground level heat radiation limits to
be exceeded, then a retrofit application is not viable.
Many flare systems are currently operated in conjunction with baseload
gas recovery systems. These systems recover and compress the waste VOC
for use as a feedstock in other processes or as fuel. When baseload gas
recovery systems are applied, the flare is used in a backup capacity and for
emergency releases. Depending on the quantity of usable VOC that can be
recovered, there can be a considerable economic advantage over operation of
a flare alone.
Streams containing high concentrations of haiogenated or sulfur contain-
ing compounds are not usually flared due to corrosion of the flare tip or
formation of secondary pollutants (such as SO?). If these vent types are to
be controlled by combustion, thermal incineration, followed by scrubbing to
remove the acid gases, is the preferred method.[3]
7-7
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7.1.3 Performance
This section discusses the parameters that affect flare VOC destruction effi-
ciency and presents the specifications that must be followed when flares are
used to comply with EPA air emission standards.
7.1.3.1 Factors Affecting Efficiency
The major factors affecting flare_combustion efficiency are vent gas flamma-
bility, auto-ignition temperature, heating value (Btu/scf), density, and flame
zone mixing.
The flammability limits of the flared gases influence ignition stability and
flame extinction. The flammability limits are defined as the stoichiometric
composition limits (maximum and minimum) of an oxygen-fuel mixture that
will burn indefinitely at given conditions of temperature and pressure without
further ignition. In other words, gases must be within their flammability
limits to burn. When flammability limits are narrow, the interior of the flame
may have insufficient air for the mixture to burn. Fuels, such as hydrogen,
with wide limits of flammability are therefore easier to combust.
For most vent streams, the heating value also affects flame stability, emis-
sions, and flame structure. A lower heating value produces a cooler flame that
does not favor combustion kinetics and is also more easily extinguished. The
lower flame temperature also reduces buoyant forces, which reduces mixing.
The density of the vent stream also affects the structure and stability
of the flame through the effect on buoyancy and mixing. By design, the
velocity in many flares is very low; therefore, most of the flame structure is
developed through buoyant forces as a result of combustion. Lighter gases
therefore tend to burn better. In addition to burner tip design, the density
also directly affects the minimum purge gas required to prevent flashback,
with lighter gases requiring more purge. (51
Poor mixing at the flare tip is the primary cause of flare smoking when
burning a given material. Streams with high carbon-to-hydrogen mole ratio
(greater than 0.35) have a greater tendency to smoke and require better
mixing for smokeless flaring.[3j For this reason one generic steam-to-vent gas
ratio is not necessarily appropriate for all vent streams. The required steam
7-8
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rate is dependent on the carbon to hydrogen ratio of the gas bdng flared. A
high ratio requires more steam to prevent a smoking flare.
7.1.3.2 Flare Specifications
At too high an exit velocity, the flame can lift off the tip and flame out, while
at too low a velocity, it can burn back into the tip or down the sides of the
stack.
The EPA requirements -foi flares used to comply with EPA air emission
standards are specified in 40 CFR Section 60.18. The requirements are for
steam-assisted, air-assisted, and non-assisted flares. Requirements for steam-
assisted, elevated flares state that the flare shall be designed for and operated
with:
• an exit velocity at the flare tip of less than 60 ft/sec for 300 Btu/scf
gas streams and less than 400 ft/sec for > 1,000 Btu/scf gas streams.
For gas streams between 300-1,000 Btu/scf the maximum permitted
velocity (Vmax, in ft/sec) is determined by the following equation:
Bv + 1214
= — t7-1)
where B,, is the net heating value in Btu/scf.
• no visible emissions. A five-minute exception period is allowed during
any two consecutive hours.
• a flame present at all times when emissions may be vented. The pres-
ence of a pilot flame shall be monitored using a thermocouple or equiv-
alent device.
• the net heating value of the gas being combusted being 300 Btu/scf or
greater.
In addition, owners or operators must monitor to ensure that flares are
operated and maintained in conformance with their design.
7-9
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7,,2 Process Description
The elements of an elevated steam-assisted flare generally consist of gas vent
collection piping, utilities (fuel, steam, and air), piping from the base up,
knock-out drum, liquid seal, flare stack, gas seal, burner tip, pilot burners,
steam jets, ignition system, and controls. Figure 7.1 is a diagram of a steam-
assisted elevated smokeless flare system showing the usual components that
are included.
7.2.1 Gas Transport Piping
x*
Process vent streams are sent from the facility release point to the flare
location through the gas collection header. The piping (generally schedule
40 carbon steel) is designed to minimize pressure drop. Ducting is not used as
it is more prone to air leaks. Valving should be kept to an absolute minimum
and should be "car-sealed" (sealed) open. Pipe layout is designed to avoid
any potential dead legs and liquid traps. The piping is equipped for purging
so that explosive mixtures do not occur in the flare system either on start-up
or during operation.
7.2.2 Knock-out Drum
Liquids that may be in the vent stream gas or that may condense out in
the collection header and transfer lines are removed by a knock-out drum.
(See Figure 7.2.) The knock-out or disentrainment drum is typically either
a horizontal or vertical vessel located at or close to the base of the flare, or
a vertical vessel located inside the base of the flare stack. Liquid in the vent
stream can extinguish the flame or cause irregular combustion and smoking.
In addition, flaring liquids can generate a spray of burning chemicals that
could reach ground level and create a safety hazard. For a flare system
designed to handle emergency process upsets this drum must be sized for
worst-case conditions (e.g., loss of cooling water or total unit depressuring)
and is usually quite large. For a flare system devoted only to vent stream
VOC control, the sizing of the drum is based primarily on vent gas flow rate
with consideration given to liquid entrainment. '
7-10
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StMinNozzlM
(9)
Riot Burrram
(7)
G«« Barrier
(6)
Helps Prevent Flash Sack
Rare Stack
(5)
Gaa Collection Header
Vert Stream J^~
Knock-out
Drum -•
(2)
Dram
Liquid
Seal -
(3)
<
1— Steam Line
Ignition
Device
Air Line
> Ga* Line
u
Figure 7.1: Steam-Assisted Elevated Flare Syst
em
7-11
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CondOTMd/Entrairwd
Uqwa
• To Slang*
Figure 7.2: Typical Vertical Knock-out Drum
7-12
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7.2.3 Liquid Seal
Process vent streams are usually passed through a liquid seal before going to
the flare stack. The liquid seal can be downstream of the knock-out drtim or
incorporated into the same vessel. This prevents possible flame flashbacks,
caused when air is inadvertently introduced into the flare system and the
flame front pulls down into the stack. The liquid seal also serves to maintain
a positive pressure on the upstream system and acts as a mechanical damper
on any explosive shock wave in the flare stack.[5] Other devices, such as flame
arresters and check valves, may sometimes replace a liquid seal or be used in
conjunction with it. Purge gas (as discussed in Section 7.3.4) also helps to
prevent flashback in the flare stack caused by low vent gas flow.
7.2.4 Flare Stack
For safety reasons a stack is used to elevate the flare. The flare must be
located so that it does not present a hazard to surrounding personnel and
facilities. Elevated flares can be self-supported (free-standing), guyed, or
structurally supported by a derrick. Examples of these three types of ele-
vated flares are shown in Figures 7.3, 7.4, and 7.5 for self-supported, derrick-
supported, and guy-supported flares, respectively. Self-supporting flares are
generally used for lower flare tower heights (30-100 feet) but can be designed
for up to 250 feet. Guy towers are designed for over 300 feet, while derrick
towers are designed for above 200 feet.[4, 6, 7, 8, 9, 10]
Free-standing flares provide ideal structural support. However, for verv
high units the costs increase rapidly. In addition, the foundation required
and nature of the soil must be considered.
Derrick-supported flares can be built as high as required since the system
load is spread over the derrick structure. This design provides for differential
expansion between the stack, piping, and derrick. Derrick-supported flares
are the most expensive design for a given flare height.
The guy-supported flare is the simplest of ail the support methods. How-
ever, a considerable amount of land is required since the guy wires are widely
spread apart. A rule of thumb for space required to erect a guy-supported
flare is a circle on the ground with a radius equal to the height of the flare
stack.{6]
7-13
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rrr
1
I
A
Figure 7.3: Self-Supported Elevated Flare
7-14
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Figure 7.4: Derrick-Supported Elevated Flare
7-15
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Figure 7.5: Guy-Supported Elevated Flare
7-16
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7.2.5 Gas Seal
Air may tend to flow back into a flare stack due to wind or the thermal
contraction of stack gaaea and create an explosion potential. To prevent this,
a gas seal is typically installed in the flare stack. One type of gas seal (also
referred to as a flare seal, stack seal, labyrinth seal, or gas barrier) is located
below the flare tip to impede the flow of air back into the flare gas network.
There are also "seals" which act as orifices in the top of the stack to reduce
the purge gas volume for a given velocity and also interfere with the passage
of air down the stack from the upper rim. These are known by the names
"internal gas seal, fluidic-seal, and arrester seal".[5] These seals are usually
proprietary in design, and their presence reduces the operating purge gas
requirements.
7.2.6 Burner Tip
The burner tip, or flare tip, is designed to give environmentally acceptable
combustion of the vent gas over the flare system's capacity range. The burner
tips are normally proprietary in design. Consideration is given to flame stabil-
ity, ignition reliabi'ity. end noise suppression. The maximum and minimum
capacity of a f.are to burn a flared gas with a stable flame (not necessarily
smokeless) is a function of tip design. Flame stability can be enhanced by
flame holder retention devices incorporated in the flare tip inner circumfer-
ence. Burner tips with modern flame holder designs can have a stable flame
over a flare gas exit velocity range of 1 to 600 ft/sec.(2] The actual maximum
capacity of a flare tip is usually limited by the vent stream pressure avail-
able to overcome the system pressure drop. Elevated flares diameters are
normally sized to provide vapor velocities at maximum throughput of about
50 percent of the sonic velocity of the gas subject to the constraints of CFR
7.2.7 Pilot Burners
EPA regulations require the presence of a continuous flame. Reliable ignition
is obtained by continuous pilot burners designed for stability and positioned
around the outer perimeter of the flare tip. The pilot burners are ignited
7-17
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by an ignition source system, which can be designed for either manual or
automatic actuation. Automatic systems are generally activated by a flame
detection device using either a thermocouple, an infra-red sensor or, more
rarely, (for ground flare applications) an ultra-violet sensor.[4]
7.2.8 Steam Jets
A diffusion flame receives its combustion oxygen by diffusion of air into the
flame from the surrounding atmosphere. The high volume of fuel flow in
a flare may require more combustion air at a faster rate than simple gas
diffusion can supply. High velocity steam injection nozzles, positioned around
the outer perimeter of the flare tip, increase gas turbulence in the flame
boundary zones, drawing in more combustion air and improving combustion
efficiency. For the larger flares, steam can also be injected concentrically into
the flare tip.
The injection of steam into a flare flame can produce other results in
addition to air entrainment and turbulence. Three mechanisms in which
steam reduces smoke formation have been presented.[1] Briefly, one theory
suggests that steam separates the hydrocarbon molecule, thereby minimizing
polymerization, and forms oxygen compounds that burn at a reduced rate
and temperature not conducive to cracking and polymerization. Another
theory claims that water vapor reacts with the carbon particles to form CO,
C02, and H2, thereby removing the carbon before it cools and forms smoke.
An additional effect of the steam is to reduce the temperature in the core
of the flame and suppress thermal cracking.[5| The physical limitation on
the quantity of steam that can be delivered and injected into the flare flame
determines the smokeless capacity of the flare. Smokeless capacity refers to
the volume of gas that can be combusted in a flare without smoke generation.
The smokeless capacity is usually less than the stable flame capacity of the
burner tip.
Significant disadvantages of steam usage are the increased noise and cost.
Steam aggravates the flare noise problem by producing high-frequency jet
noise. The jet noise can be reduced by the use of small multiple steam
jets and, if necessary, by acoustical shrouding. Steam injection is usually
controlled manually with the operator observing the flare (either directly or
on a television monitor) and adding steam as required to maintain smokeless
operation. To optimize steam usage infrared sensors are available that sense
7-18
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flare flame characteristics and adjust the steam flow rate automatically to
maintain smokeless operation. Automatic control, based on flare gas flow
and flame radiation, gives a faster response to the need for steam and a
better adjustment of the quantity required. If a manual system is used, steam
metering should be installed to significantly increase operator awareness and
reduce steam consumption.
7.2.9 Controls
Flare system control can be completely automated or completely manual.
Components of a flare system which can be controlled automatically include
the auxiliary gas, steam injection, and the ignition system. Fuel gas con-
sumption can be minimized by continuously measuring the vent gas flow
rate and heat content (Btu/scf) and automatically adjusting the amount of
auxiliary fuel to maintain the required minimum of 300 Btu/scf for steam-
assisted flares. Steam consumption can likewise be minimized by controlling
flow based on vent gas flow rate. Steam flow can also be controlled using vi-
sual smoke monitors. Automatic ignition panels sense the presence of a flame
with either visual or thermal sensors and reignite the pilots when flameouts
occur.
7.3 Design Procedures
Flare design is influenced by several factors, including the availability oi"
space, the characteristics of the flare gas (namely composition, quantity, and
pressure level) and occupational concerns. The sizing of flares requires deter-
mination of the required flare tip diameter and height. The emphasis of this
section will be to size a steam-assisted elevated flare for a. given application.
7.3.1 Auxiliary Fuel Requirement
The flare tip diameter is a function of the vent gas flow rate plus the auxiliary
fuel and purge gas flow rates. The purge gas flow rate is very small relative to
the vent gas and fuel flow rates, so it may be ignored when determining the
7-19
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tip diameter. The flow rate of the auxiliary fuel, if required, w significant,
and must be calculated before the tip diameter can be computed.
Some flares are provided with auxiliary fuel to combust hydrocarbon va-
pors when a lean flare gas stream falls below the flammability range or heating
value necessary to sustain a stable flame. The amount of fuel required, F,
is calculated based on maintaining the vent gas stream net heating value at
the minimum of 300 Btu/scf required by rules defined in the Federal Register
(see next section):
Q Bv + F Bf = (Q + F)(300 Btu/scf) (7.2)
where
Q = the vent stream flow rate, scfm
Bv and B/ are the Btu/scf of the vent stream and fuel, respectively.
Rearranging gives:
The annual auxiliary fuel requirement, Fa, is calculated by:
Fa (Mscf/yr) = (F scfm)(60 min/hr)(8760 hr/yr)
= 526^ (7.4)
Typical natural gas has a net heating value of about 1,000 Btu/scf. Auto-
matic control of the auxiliary fuel is ideal for processes with large fluctuations
in VOC compositions. These flares are used for the disposal of such streams
as sulfur tail gases and ammonia waste gases, as well as any low Btu vent
streams. [2]
7.3.2 Flare Tip Diameter
Flare tip diameter is generally sized on a velocity basis, although pressure
drop must also be checked. Flare tip sizing for flares used to comply with EPA
air emission standards is governed by rules defined in the Federal Register
(see 40 CFR 60.18). To comply with these requirements, the maximum
velocity of a steam-assisted elevated flare is determined as follows:
7-20
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Net Heating Value of
Vent Stream Maximum Velocity
Bv (Btu/scf) Fmax (ft/sec)
300 60
300 - 1,000 loglo (Vmax) = (Bv + l,214)/852
> 1,000 400
By determining the maximum allowed velocity, Fmax (ft/sec), £i.nd know-
ing the total volumetric flow rate, Qtot (acfm), including vent stream and
auxiliary fuel gas, a minimum flare tip diameter, Z?min (in), can be calcu-
lated. It is standard practice to size the flare so that the design velocity of
flow rate is 80 percent of Vmax, i.e.:
Qtnt
* 60 (sec/min)
0.8 Vmax .
/ , is the calculated diameter, D = Z>mjn, rounded
up to the next commercially available size. The minimum flare size is 1 inch;
larger sizes are available in 2-inch increments from 2 to 24 inches ind in 6-
inch increments above 24 inches. The maximum size commercially available
is 90 inches.[5]
A pressure drop calculation is required at this point to ensure that the
vent stream has sufficient pressure to overcome the pressure drop occurring
through the flare system at maximum flow conditions. The pressure drop
calculation is site specific but must take into account losses through the
collection header and piping, the knock-out drum, the liquid seal, the flare
stack, the gas seal, and finally the flare tip. Piping size should be assumed
equal to the flare tip diameter. Schedule 40 carbon steel pipe is typically
used. If sufficient pressure is not available, the economics of either a larger
flare system (pressure drop is inversely proportional to the pipe dianeter) or
a mover such as a fan or compressor must be weighed. (Refer to Section 7.3.8
for typical pressure drop relationships.)
7-21
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7.3.3 Flare Height
The height of a flare is determined based on the ground level limitations of
thermal radiation intensity, luminosity, noise, height of surrounding struc-
tures, and the dispersion of the exhaust gases. In addition, consideration
must also be given for plume dispersion in case of possible emission ignition
failure. Industrial flares are normally sized for a maximum heat intensity of
1,500-2,000 Btu/hr-ft2 when flaring at their maximum design rates. [1, 2] At
this heat intensity level, workers can remain in the area of the flare for a
limited period only. If, however, operating personnel are required to remain
in the unit area performing their duties, the recommended design flare ra-
diation level excluding solar radiation is 500 Btu/hr-ft2. [1] The intensity of
solar radiation is in the range of 250-330 Btu/hr-ft2. [1] Flare height may also
be determined by the need to safely disperse the vent gas in case of flameout.
The height in these cases would be based on dispersion modeling for the
particular installation conditions and is not addressed here. The minimum
flare height normally used is 30 feet. [5] Equation (7.6) by Hajek and Ludwig
may be used to determine the minimum distance, £, required from the center
of the flare flame and a point of exposure where thermal radiation must be
limited.fl]
where
r = fraction of heat intensity transmitted
/ = fraction of heat radiated
R = net heat release (Btu/hr)
K = allowable radiation (500 Btu/hr-ft2)
The conservative design approach used here ignores wind effects and cal-
culates the distance assuming the center of radiation is at the base of the
flame (at the flare tip), not in the center. It is also assumed that the lo-
cation where thermal radiation must be limited is at the base of the flare.
Therefore, the distance, £, is equal to the required flare stack height (which
is a minimum of 30 feet). The / factor allows for the fact that not all the
heat released in a flame can be released as radiation. Heat transfer is prop-
agated through three mechanisms: conduction, convection, and radiation.
Thermal radiation may be either absorbed, reflected, or transmitted. Since
the atmosphere is not a perfect vacuum, a fraction of the heat radiated is
not transmitted due to atmospheric absorption (humidity, particulate mat-
7-22
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ter). For estimating purposes, however, assume all of the heat radiated is
transmitted (i.e., T = 1). The following is a summary of heat radiated from
various gaseous diffusion flames:[1]
Gas Flare Tip Diameter (in) Fraction of Heat Radiated (/)
Hydrogen <1 .10
1.6 .11
3.3 .16
8.0 .15
. 16.0 .17
Butane
Methane
Natural gas
<1
1.6
3.3
8.0
16.0
<1
1.6
3.3
8.0
16.0
.29
.29
.29
.28
.30
.16
.16
.15
.19
.23
In general, the fraction of heat radiated increases as the staick diameter
increases. If stream-specific data are not available, a design basis of / = 0.2
will give conservative results.[4] The heat release, R, is calculated from the
flare gas flow rate, W, and the net heating value, B,n as follows:
R (Btu/hr) = (W lb/hr)(S,, Btu/lb) (7.7)
7.3.4 Purge Gas Requirement
The total volumetric flow to the flame must be carefully controlled to pre-
vent low flow flashback problems and to avoid flame instability. Purge gas,
typically natural gas, N2, or C02, is used to maintain a minimum required
positive flow through the system. If there is a possibility of air in the flare
7-23
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manifold, N2, another inert gas, or a flammable gas must be used to prevent
the formation of an explosive mixture in the flare system. To ensure a posi-
tive flow through all flare components, purge gas injection should be at the
farthest upstream point in the flare transport .piping.
The minimum continuous purge gas required is determined by the design
of the stack seals, which are usually proprietary devices. Modern labyrinth
and internal gas seals are stated to require a gas velocity of 0.001 to 0.04
ft/sec (at standard conditions). [6, 7, 8, 9, 10] Using the conservative value of
0.04 ft/sec and knowing the flare diameter (in), the annual purge gas volume,
Fpuj can be calculated:
F^ (Mscf/yr) = (0.04 ft/sec) — f-^ (3,6^0 sec/hr)(8,760 hr/yr)
I 144 it I
= 6.88D2 (Mscf/yr) (7.8)
There is another minimum flare tip velocity for operation without burn lock
or inatability. This minimum velocity is dependent on both gas composition
and diameter and can range from insignificant amounts on small flares to 0.5
ft/sec on greater than 60-inch diameter units. [5]
Purge gas is also required to clear the system of air before startup, and
to prevent a vacuum from pulling air back into the system after a hot gas
discharge is flared. (The cooling of gases within the flare system can create
a vacuum.) The purge gas consumption from these uses is assumed to be
minor.
7.3.5 Pilot Gas Requirement
The number of pilot burners required depends on flare size and, possibly, on
flare gas composition and wind conditions. Pilot gas usage is a function of
the number of pilot burners required to ensure positive ignition of the flared
gas, of the design of the pilots, and of the mode of operation. The average
pilot gas consumption based on an energy-efficient model is 70 scf/hr (of
typical 1000 Btu per scf gas) per pilot burner.[6, 7, 8, 9, 10] The number of
pilot burners, N, based on flare size is:(6, 7, 8, 9, 10]
7-24
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Flare Tip Diameter (in) Number of Pilot Burners (AT)
TTo ~ I
12-24 2
30-60 3
>60 4
The annual pilot gas consumption, Fpi, is calculated by:
Fpi (Mscf/yr) = (70 scf/hr)(^V)(8,760 hr/yr)
V (7.9)
7.3.6 Steam Requirement
The steam requirement depends on the composition of the vent gas being
flared, the steam velocity from the injection nozzle, and the flare tip diameter.
Although some gases can be flared smokelessly without any steam, typically
0.01 to 0.6 pound of steam per pound of flare gas is required. (6, 7, 8, 9, 10] The
ratio is usually estimated from the molecular weight of the gas, the carbon-
to-hydrogen ratio of the gas, or whether the gas is saturated or unsaturated.
For example, olefins, such as propylene, require higher steam ratios than
would paraffin hydrocarbons to burn smokelessly. [2]
In any event, if a proprietary smokeless flare is purchased, the manufac-
turer should be consulted about the minimum necessary steam rate. A small
diameter flare tip (less than 24 inches) can use steam more effectively than a
large diameter tip to mix air into the flame and promote turbulence. [2 j For
a typical refinery, the average steam requirement is typically 0.25 Ib/lb, with
this number increasing to 0.5 Ib/lb in chemical plants where large quantities
of unsaturated hydrocarbons are flared. (10)
For general consideration, the quantity of steam required, 5, cz.n be as-
sumed to be 0.4 pounds of steam per pound of flare gas, W. Using a 0.4
ratio, the amount of steam required is:
5(lbs/yr) = QA(W lb/hr)(8, 760 hr/yr)
= 3,500(11' Ibs/hr) (7.10)
Operating a flare at too high a steam-to-gas ratio is not only costly,
but also results in a lower combustion efficiency and a noise nuisarce. The
7-25
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capacity of a steam-assisted flare to burn smokelessly may be limited by the
quantity of steam that is available.
7.3.7 Knock-out Drum
As explained previously, the knock-out drum is used to remove any liquids
that may be in the vent stream. Two types of drums are used: horizontal and
vertical. The economics of vessel design influences the choice between a hori-
zontal and a vertical drum. When a large liquid storage vessel is required and
the vapor flow is high, a horizontal drum is usually more economical. Ver-
tical separators are used when there is small liquid load, limited plot space,
or where ease of level control is desired. It is assumed here that the drum
is not sized for emergency releases and that liquid flow is minimal. Flares
designed to control continuous vent streams generally have vertical knock-
out drums, whereas emergency flares typically have horizontal vessels. The
procedure described below applies to vertical drums exclusively. A typical
vertical knock-out drum is presented in Figure 7.2.
Liquid particles will separate when the residence time of the vapor is
greater than the time required to travel the available vertical height at the
dropout velocity of the liquid particles, i.e., the velocity is less than the
dropout velocity. In addition, the vertical gas velocity must be sufficiently low
to permit the liquid droplets to fall. Since flares are designed to handle small-
sized liquid droplets, the allowable vertical velocity is based on separating
droplets from 300 to 600 micrometers in diameter. [1] The dropout velocity, U,
of a particle in a stream, or the maximum design vapor velocity, is calculated
as followsrf 11]
U (ft/sec) = cJ^-^^ (7.11)
V Pv
where
G = design vapor velocity factor
p\ and pv = liquid and vapor densities, lb/ft:I
Note that in most cases,
P\ ~ P<- P
r
The design vapor velocity factor, G, ranges from 0.15 to 0.25 for vertical
gravity separators at 85% of flooding. (11)
7-26
-------
Once the maximum design vapor velocity has been determined the mini-
mum vessel cross-sectional area, A, can be calculated by:
A fft2% _ Q* ft3/™
1 ;
(60 sec/min)(tf ft/sec)
where 3 condition.
So for purposes of flare knock-out drum sizing:
h (in) = Zd (7.15)
7-27
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7.3.8 Gas Mover System
The total system pressure drop is a function of the available pressure of
the vent stream, the design of the various system components, and the flare
gas flow rate. The estimation of actual pressure drop requirements involves
complex calculations based on the specific system's vent gas properties and
equipment used. For the purposes of this section, however, approximate
values can be used. The design pressure drop through the flare tip can
range from « 0.1 to 2 psi with the following approximate pressure drop
relationships:[5]
Gas seal: 1 to 3 times flare tip pressure drop
Stack: 0.25 to 2 times flare tip pressure drop
Liquid seal and Knock- 1 to 1.5 times flare tip pressure drop plus
out drum: pressure drop due to liquid depth in the
seal, which is normally 0.2 to 1.5 psi.
Gas collection system: calculated based on diameter, length, and
flow. System is sized by designer to utilize
the pressure drop available and still leave
a pressure at the stack base of between 2
and 10 psi.
Typical total system pressure drop ranges from about 1 to 25 psi.[5]
7.4 Estimating Total Capital Investment
The capital costs of a flare system are presented in this section and are based
on the design/sizing procedures discussed in Section 7.3. The costs presented
are in March 1990 dollars.
Total capital investment, TCI/includes the equipment costs, EC, for the
flare itself, the cost of auxiliary equipment, the cost of taxes, freight, and
instrumentation, and all direct and indirect installation costs.
The capital cost of flares depends on the degree of sophistication desired
(t.e., manual vs automatic control) and the number of appurtenances se-
lected, such as knock-out drums, seals, controls, ladders, and platforms. The
7-28
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basic support structure of the flare, the size and height, and the auxiliary
equipment are the controlling factors in the cost of the flare. Tli.e capital
investment will also depend on the availability of utilities such as steam,
natural gas, and instrument air.
The total capital investment is a battery limit cost estimate and does not
include the provisions for bringing utilities, services, or roads to the; site, the
backup facilities, the land, the research and development required, or the
process piping and instrumentation interconnections that may be required
in the process generating the waste gas. These costs are based on a new
plant installation; no retrofit cost considerations such as demolition, crowded
construction working conditions, scheduling construction with production
activities, and long interconnecting piping are included. These factors are so
site-specific that no attempt has been made to provide their costs.
7.4.1 Equipment Costs
Flare vendors were asked to provide budget estimates for the spectrum of
commercial flare sizes. These quotes [6, 7, 8, 9, 10] were used to develop
the equipment cost correlations for flare units, while the cost equations for
the auxiliary equipment were based on references [12] and [13] (knock-out
drums) and [14] and [15] (piping). The expected accuracy of these costs is
± 30% (i.e., "study" estimates). Keeping in mind the height restrictions
discussed in Section 7.2.4, these cost correlations apply to flare tip diameters
ranging from 1 to 60 inches and stack heights ranging from 30 to 500 feet.
The standard construction material is carbon steel except when it is standard
practice to use other materials, as is the case with burner tips.
The flare costs, CV, presented in Equations 7.16 through 7.18 are calcu-
lated as a function of stack height, L (ft) (30 ft minimum), and tip diameter,
D (in), and are based on support type as follows:
•
Self Support Group:
CV (S) = (78.0 + 9.140 + 0.749/i)-' (7.16)
Guy Support Group:
CV (S) = (103 4- 8.681? 4- 0.4701)- (7.17)
7-29
-------
Derrick Support Group:
C,, ($) = (76.4 + 2.721? + 1.64X)2 (7.18)
The equations are least-squares regression of cost data provided by differ-
ent vendors. It must be kept in mind that even for a given flare technology
(i.e., elevated, steam-assisted), design and manufacturing procedures vary
from vendor to vendor, so that costs may vary. Once a study estimate is
completed, it is recommended that several vendors be solicited for more de-
tailed cost estimates.
Each of these costs includes the flare tower (stack) and support, burner
tip, pilots, utility (steam, natural gas) piping from base, utility metering and
control, liquid seal, gas seal, and galvanized caged ladders and platforms as
required. Costs are based on carbon steel construction, except for the upper
four feet and burner tip, which are based on 310 stainless steel.
The gas collection header and transfer line requirements are very site
specific and depend on the process facility where the emission is generated
and on where the flare is located. For the purposes of estimating capital cost
it is assumed that the transfer line will be the same diameter as the flare
tip[6] and will be 100 feet long. Most installations will require much more
extensive piping, so 100 feet is considered a minimum.
The costs for vent stream piping, Cp, are presented separately in Equation
7.19 or 7.20 and are a function of pipe, or flare, diameter, ZJ.fiS]
CP ($) = 127D'-21 (where 1"< D <24") (7.19)
•Cp ($) = 139£>'-u7 (where 30"< D <60") (7.20)
The costs, C/>, include straight, Schedule 40, carbon steel pipe only, are based
on 100 feet of piping, and are directly proportional to the distance required.
The costs for a knock-out drum, CV, are presented separately in Equation
7.21 and are a function of drum diameter, d (in), and height, h (in).[12, 13]
CK (S) = 14.2[«f t (h + 0.812d)]"'7:ir (7.21)
where t is the vessel thickness, in inches, determined based on the diameter.
Flare system equipment cost, EC, is the total of the calculated flare,
knock-out drum, and piping costs.
EC ($) = CV + C,, + Cr (7.22)
7-30
-------
Purchased equipment costs, PEC, is equal to equipment cost, EC, plus factors
for ancillary instrumentation (i.e., control room instruments) (.10), sales
taxes (0.03), and freight (0.05) or,
PEC ($) = EC (1 + 0.10 + 0.03 4- 0.05) = 1.18 EC (7.23)
7.4.2 Installation Costs
The total capital investment, TCI, is obtained by multiplying the purchased
equipment cost, PEC, by an installation factor of 1.92.
' TCI ($) = 1.92 PEC (7.24)
These coats were determined based on the factors in Table 7.1. Thesie factors
encompass direct and indirect installation costs. Direct installation costs
cover foundations and supports, equipment" handling and erection, piping,
insulation, painting, and electrical. Indirect installation costs cover engineer-
ing, construction and field expenses, contractor fees, start-up, performance
testing, and contingencies. Depending on the site conditions, the installa-
tion costs for a given flare could deviate significantly from costs generated
by these average factors. Vatavuk and Neveril provide some guidelines for
adjusting the average installation factors to account for other-than-average
installation conditions.[l6]
7.5 Estimating Total Annual Costs
The total annual cost, TAG, is the sum of the direct and indirect annual costs.
The bases used in calculating annual cost factors are given in Table 7.2.
7.5.1 Direct Annual Costs
Direct annual costs include labor (operating and supervisory), maintenance
(labor and materials), natural gas, steam, and electricity. Unless the flare is
to be dedicated to one vent stream and specific on-line operating factors are
known, costs should be calculated based on a continuous operation of 8,760
7-31
-------
Table 7.1: Capital Cost Factors for Flare Systems
Cost Item Factor
Direct Costs
Purchased equipment costs
Flare system, EC As estimated, A
Instrumentation 0.10 A
Sales taxes . 0.03 A
Freight * 0.05 A
Purchased equipment cost, PEC B = 1.18 A
Direct installation costs
Foundations & supports 0.12 B
Handling & erection 0.40 B
Electrical 0.01 B
Piping 0.02 B
Insulation 0.01 B
Painting 0.01 B
Direct installation costs 0.57 B
Site preparation As required, SP
Buildings As required, Bldg.
Total Direct Costs, DC 1.57 B + SP 4- Bldg.
Indirect Costs (installation)
Engineering 0.10 B
Construction and field expenses 0.10 B
Contractor fees 0.10 B
Start-up 0.01 B
Performance test 0_Q1 g
Contingencies 0.03 B
Total Indirect Costs. 1C 0.35 3
Total Capital Investment = DC •*- 1C 1.92 B - SP J- Bldg.
r-32
-------
Table 7.2: Suggested Annual Cost Factors for Flare Systems
Cost Item
Direct Annual Costs, DC
Operating labor(3] __
Operator
Supervisor
Operating materials
Maintenance
Labor
Material
Utilities
Electricity
Purge gas
Pilot gas
Auxiliary fuel
Steam
Indirect Annual Costs, 1C
Overhead
Administrative charges
Property tax
Insurance
Capital recovery"
Total Annual Cost
Factor
630 man-hours/year
15% of operator
1/2 hour per shift
100% of maintenance labor
All utilities equal to:
(consumption rate) x
(hours/yr) x (unit cost)
60% of total labor and material costs
2% of Total Capital Investment
1% of Total Capital Investment
1% of Total Capital Investment
0.1315 x Total Capital Investment
Sum of Direct and Indirect Annuail Cbsts
see Chapter 2.
7-33
4
-------
hr/yr and expressed on an annual basis. Flares serving multiple process units
typically run continuously for several years between maintenance shutdowns.
Operating labor is estimated at 630 hours annually.[3] A completely man-
ual system could easily require 1,000 hours. A standard supervision ratio of
0.15 should be assumed.
Maintenance labor is estimated at 0.5 hours per 8-hour shift. Maintenance
materials costs are assumed to equal maintenance labor costs. Flare utility
costs include natural gas, steam, and electricity.
Flare systems can use natural gas in three ways: in pilot burners that fire
natural gas, in combusting low Btu vent streams that require natural gas as
auxiliary fuel, and as purge gas. The total natural gas cost, C/, to operate a
flare system includes pilot, Cpl, auxiliary fuel, Ca, and purge costs, Cpll:
Cj ($/yr) = C,n + Ca + Cpu (7.25)
where, Cp, is equal to the annual volume of pilot gas, F,,t, multiplied by the
cost per scf, i.e.:
CP< ($/yr) = (Fp, scf/yr)(S/scf) (7.26)
Cn and Cj», are similarly calculated.
Steam cost (C,) to eliminate smoking is equal to the annual steam con-
sumption 8,760 S multiplied by the cost per Ib, i.e.:
C, ($/yr) = (8,760 hr/yr)(S lb/hr)($/lb) (7.27)
The use of steam as a smoke suppressant can represent as much as 90% or
more of the total direct annual costs.
7.5.2 Indirect Annual Costs
The indirect (fixed) annual costs include overhead, capital recovery, admin-
istrative (G&A) charges, property taxes, and insurance. Suggested indirect
annual cost factors are presented in Table 7.2.
Overhead is calculated as 60% of the total labor (operating, maintenance,
and supervisory) and maintenance material costs. Overhead cost is discussed
in Chapter 2 of this Manual.
7-34
-------
The system capital recovery cost, CRC, is based on an estimated 15-year
equipment life. (See Chapter 2 of this Manual for a thorough discussion of
the capital recovery cost and the variables that determine it.) For a 15-year
life and an interest rate of 10%, the capital recovery factor is 0.1315. The
system capital recovery cost is the product of the system capital recovery
factor, CRF, and the total capital investment, TCI, or:
CRC ($/yr) = CRF x TCI = 0.1315 x TCI (7.28)
As shown in Table 7.2, G& A, taxes, and insurance can be estimated at 2%,
1%, and 1% of the total capital investment, TCI, respectively.
7.6 Example Problem
The example problem described in this section shows how to apply the flare
sizing and costing procedures to the control of a vent stream associated with
the distillation manufacturing of methanol.
7.6.1 Required Information for Design
The first step in the design procedure is to determine the specifications of
the vent gas to be processed. The minimum information required to size a
flare system for estimating costs are the vent stream:
Volumetric or mass flow rate
Heating value or chemical composition
Temperature
System pressure
Vapor and liquid densities
In addition the following are needed to calculate direct annual costs':
Labor costs
Fuel costs
Steam costs
Vent stream parameters and cost data to be used in this example problem
are listed in Table 7.3.
7-35
-------
Table 7.3: Example Problem Data
Vent Stream Parameters"
Flow rate
Heat content
System pressure
Temperature
Liquid density[17]
Vapor density (17)
63.4 acfm"
^399.3 Ib/hr
449 Btu/scf6
10 psigc
90 °F
49.60 lb/ft3
0.08446 lb/ft:j
Cost Data (March 1990)[18, 19]
Operating hours
Natural gas
Steam
Operating labor
Maintenance labor
8,760
3.03
4.65
15.64
17.21
hrs/yr
S/1000 scf
S/1000 Ibs
$/hr
S/hr
"Measured at flare tip. Flow rate has been adjusted to account
for drop in pressure from 10 psig at source to 1 psig at flare tip.
^Standard conditions: 77°F, 1 atmosphere.
Treasure at source (gas collection point). Pressure at flare tip
is lower: 1 psig.
7-36
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7.6.2 Capital Equipment
The first objective is to properly size a steam-assisted flare system to effec-
tively destroy 98% of the VOC (methanol) in the vent gas stream. Using the
vent stream parameters and the design procedures outlined in Section 7.3,
flare and knock-out drum heights and diameters can be determined. Once
equipment has been specified, the capital costs can be determined from equa-
tions presented in Section 7.4.1.
7.6.2.1 Equipment Design
The first step in flare sizing is determining the appropriate flare tip diameter.
Knowing the net (lower) heating value of the vent stream, the maximum
allowed velocity can be calculated from the Federal Register requirements.
Since the heating value is in the range of 300 to 1,000 Btu/scf, the maximum
velocity, Vmax, is determined by Equation 7.1.
. T. 449 Btu/scf+1,214
Icg.oVmax = g^
= 1.95
so,
Vmax = 89.5 ft/sec
Because the stream heating value is above 300 Btu/scf, no auxiliary fuel is
required. Hence, Qtot equals the vent stream flow rate. Based on <;)tot and
T" the flare tip diameter can be calculated using Equation 7.5.
Anin = 1-95
= 1.95
89.5 ft/sec
= 1.64 in
The next largest commercially available standard size of 2 inches should he
selected for D.
The next parameter to determine is the required height of the flare stack.
The heat release from the flare is calculated using Equation 7.7.
R (Btu/hr) = (W lb/hr)(£T. Btu/lb)
7-37
-------
First the heat of combustion, or heating value, must be converted from
Btu/scf to Btu/lb. The vapor density of the vent stream at standard tem-
perature and pressure is 0.08446 Ib/scf. So,
449 Btu/scf
* = 0.08446 Ib/scf = 5316 Btu/lb
and,
R = (399.3 lb/hr)(5,316 Btu/lb) = 2,123,000 Btu/hr
Substituting R and appropriate values for other variables into Equation 7.6:
(1)(0.2)(2,123,QOO Btu/hr)
4^(500 Btu/hr-ft2)
= 68 ft2
gives a height of L = 8.2 ft. The smallest commercially available flare is 30
feet, so L = 30 ft.
Next the knock-out drum must be sized. Assuming a design vapor velocity
factor, (7, of 0.20, and substituting the vapor and liquid densities of methanol
into Equation 7.11 yields a maximum velocity of:
u = Gy-L_-,ft/sec
Pv
0 20i - ~ °'08446
0.08446
= 4.84 ft/sec
Given a vent gas flow rate of 63.4 scfm, the minimum vessel cross-sectional
diameter is calculated by Equation 7.12:
_ Qa
(60 sec/min)(tf ft/sec)
63.4
(60)(4.84)
= 0.218 ft-
7-38
-------
This results in a minimum vessel diameter of:
^min = 13.5vCl
= 13.5V0.218
= 6.3 inches
The selected diameter, d, rounded to the next largest 6 inches is 12 inches.
Using the rule of the height to diameter ratio of three gives a vessel height
of 36 inches, or 3 feet.
7.6.2.2 Equipment Costs
Once the required flare tip diameter and stack height have been determined
the equipment costs can be calculated. Since the height is 30 feet, the flare
will be self-supporting. The costs are determined from Equation 7.16.
CF = (78.0 + 9.14£> + Q.74,91)2
= [78.0 + 9.14(2 inches) + 0.749(30 ft)]2
= $14,100
Knock-out drum costs are determined using Equation 7.21, where t is deter-
mined from the ranges presented in Section 7.3.7. Substituting 0.215 for t:
CK = 1
= 14.2[(12)(0.25)(36 + 0.812(12))]"'7n7
= $530
Transport piping costs are determined using Equation 7.19.
CP = 127£>''21
= 127(2)''21
= S290
The total auxiliary equipment cost is the sum of the knock-out drum and
transport piping costs, or 5530 - S290 = 3820.
The total capital investment is calculated using the factors give^n in Ta-
ble 7.1. The calculations are shown in Table 7.4. Therefore:
7-39
-------
Table 7.4: Capital Costs for Flare Systems
Example Problem
Cost Item Cost
Direct Costs
Purchased equipment costs
Flare (self supporting) $14,100
Auxiliary equipment" 820
Sum = A $14,920
Instrumentation, 0.1A _ 1,490
Sales taxes, 0.03A ~ 450
Freight, 0.05A 750
Purchased equipment cost, B ,. $17,610
Direct installation costs
Foundation and supports, 0.12B 2,110
Handling & erection, 0.40B 7,040
Electrical, 0.01B 180
Piping, 0.02B 350
Insulation, 0.01B 180
Painting, 0.01B 180
Direct installation cost $10,040
Site preparation —
Facilities and buildings —
Total Direct Cost ' 527,650
Indirect Costs (installation)
Engineering, 0.10B 1,760
Construction and field expenses, 0.10B 1,760
Contractor fees, 0.10B 1,760
Start-up, 0.01B 180
Performance test, 0.01B 180
Contingencies, 0.03B " 530
Total Indirect Cost S6,170
Total Capital Investment (rounded) $33,800
"Includes costs for knock-out drum and transport piping.
7-40
-------
Purchased Equipment Cost = "B" = 1.18 x A
= 1.18 x (14,920) = $17,610
And:
Total Capital Investment (rounded) = 1.92 X B
= 1.92 x (17,610) = $33,800.
7.6.3 Operating .Requirements
Operating labor is estimated at 630 hours annually with supervisory labor at
15% of this amount. Maintenance labor is estimated at 1/2 hour per shift.
Maintenance material costs are assumed to be equal to maintenance labor
costs.
As stated above, since the heat content of the example stream is above
300 Btu/scf (i.e., 449 Btu/scf) no auxiliary fuel is needed. Natural gas
is required, however, for purge and pilot gas. Purge gas requirements are
calculated from Equation 7.8.
Fpll = 6.88£>2 = 6.88(2)2 = 27.5 Mscf/yr
Since the flare tip diameter is less than 10 inches, pilot gas requirements
are based on one pilot burner, (see Section 7.3.5) and are calculated by
Equation 7.9.
Fp, = 613N
When N = 1,
Fp, =613 Mscf/yr
Steam requirements are calculated from Equation 7.10:
5(lb/yr) = 3,500 W
Inserting the methanol mass flow rate of 399.3 Ib/hr yields:
5 = (3,500)(399.3 Ib/hr)
= 1,400 Mlb/yr
7-41
-------
7.6.4 Total Annual Costs
The sum of the direct and indirect annual costs yields a total annual cost of
$62,500. Table 7.5 shows the calculations of the direct and indirect annual
costs for the flare system as calculated from the factors in Table 7.2. Direct
costs include labor, materials, and utilities. Indirect costs are the fixed costs
allocated to the project, including capital recovery costs and such costs as
overhead, insurance, taxes, and administrative charges.
Electrical costs of a mover system (fan, blower, compressor) would have
to be included if the vent stream, pressure was not sufficient to overcome the
flare system pressure drop. In this example case, the pressure is assumed to
be adequate.
7.7 Acknowledgments
The authors gratefully acknowledge the following companies for contributing
data to this chapter:
• Flaregas Corporation (Spring Valley, NY)
• John Zink Company (Tulsa, OK)
• Kaldair Incorporated (Houston, TX)
• NAO Incorporated (Philadelphia, PA)
• Peabody Engineering Corporation (Stamford, CT)
• Piedmont HUB, Incorporated (Raleigh, NC)
7-42
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Table 7.5: Annual Costs for Flare System
Example Problem
Cost Item
Calculations
Cost
Direct Annual Costs, DC
Operating Labor
Operator
Supervisor
Operating materials
Maintenance
Labor
Material
Utilities
Electricity
Purge gas
Pilot gas
Steam
Total DC
Indirect Annual Costs, 1C
Overhead
Administrative charges
Property tax
Insurance-
Capital recovery"
Total 1C
15% of operator = 0.15 x 9,850
8,760 h 817.21
h
100% of maintenance tabor
27.5 Mscf x S3.03
613 Mscf v 83.03
yr "SiscT
1,400 x 10-' lb 8465
y * IFMb
60% of total labor and material costs:
= 0.6(9,850 -I- 1,480 + 9,420 + 9,420)
2% of Total Capital Investment = 0.02(833,800)
1% of Total Capital Investment = 0.01(833,800)
1% of Total Capital Investment = 0.01(333,800)
0.1315 x 533,800
Total Annual Cost (rounded)
59,850
1,480
9,420
9,420
80
1,860
6,510
$38,600
18,100
680
340
340
4,440
523,900
$62,500
ine capital recovery cost tactor. CRF, is a function of the flare equipment life and (lie
opportunity cost of the capital (i.e.. interest rate). For example, for a 15 vear equipment
life and a 10% interest rate, CRF = 0.1315.
7-43
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References
[1] Guide for Pressure-Relieving and Depreasurizing Systems, Refining De-
partment, API Recommended Practice 521, Second Edition, September
1982.
[2] Kalcevic, V. (IT Enviroscience), "Control Device Evaluation Flares and
the Use of Emissions as Fuels", Organic Chemical Manufacturing Vol-
ume 4i Combustion Control Devices, U.S. Environmental Protection
Agency, Research Triangle Park, NC, Publication no. EPA-450/3-30-
026, December 1980, Report 4.
[3] Reactor Processes in Synthetic Organic Chemical Manufacturing Indus-
try—Background Information for Proposed Standards, U.S. Environ-
mental Protection Agency, Office of Air Quality Planning and Stan-
dards, Research Triangle Park, NC, Preliminary Draft, EPA 450/3-90-
016a, June 1990.
[4] Letter from J. Keith McCartney (John Zink Co., Tulsa, OK) to William
M. Vatavuk (U.S. Environmental Protection Agency, Researcn Triangle
Park, NC), November 19, 1990. .
[5] Letter from David Shore (Flaregas Corp., Spring Valley, NY) to William
M. Vatavuk (U.S. Environmental Protection Agency, Research Triangle
Park, NC), October 3, 1990.
[6] Letter from Pete Tkatschenko (NAO, Inc., Philadelphia, PA) to Diana
Stone (Radian, Research Triangle Park, NC), May 2, 1990.
[7] Letter to Gary Tyler (Kaldair, Inc., Houston, TX) to Diana Stone (Ra-
dian, Research Triangle Park, NC), April 10, 1990.
[8] Letter from Zahir Bozai (Peabody Engineering Corp., Stamford, CT) to
Diana Stone (Radian, Research Triangle Park, NC), May 7, 1990.
7-44
-------
. [9] Letter from James Parker (John Zink Co., Tulsa, OK) to Diana Stone
(Radian, Research Triangle Park, NC), April 17, 1990.
(10) Letter from Nick Sanderson (Flaregas Corp., Spring Valley, NY) to Di-
ana Stone (Radian, Research Triangle Park, NC), May 2, 1990.
[11] Wu, F.H., "Drum Separator Design, A New Approach," Chemical En-
gineering, April 2, 1984, pp. 74-81.
[12] Mulct, A., "Estimate Costs of Pressure Vessels Via Correlations," Chem-
ical Engineering, October 5, 1981, pp. 145-150.
[13] Process Plant Construction Estimating Standards, Richardson Engineer-
ing Services, Inc., Volume 4, 1988 Edition.
[14] Peters, Max S. and Klaus D. Timmerhaus, Plant Design and Economics
for Chemical Engineers, Third Edition, McGraw-Hill, 1980.
[15] Cost information from Piedmont HUB, Incorporated, Raleigh, NC, Au-
gust 1990.
[16] Vatavuk, W.M., and R. Neveril, "Estimating Costs of Air Pollution
Control Systems, Part II: Factors for Estimating Capital and Operating
Costs," Chemical Engineering, Novembers, 1980, pp. 157-162.
[17] Handbook of Chemistry and Physics, 55th Edition, CRC Press, 1974-
1975.
[18] Green, G.P. and Epstein, R.K., Employment and Earnings, Department
of Labor, Bureau of Labor Statistics, Volume 37, No. 4, April 1990.
[19] Monthly Energy Review, Energy Information Administration, Office of
Energy Markets and End Use, U.S. Department of Energy, DOE-EIA-
0035(90/12), February 1990.
r-45
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Chapter 8
REFRIGERATED
CONDENSERS
Gunseli Sagun Shareef
Wiley J. Barbour
Susan K. Lynch
W. Richard Pelt
Radian Corporation
Research Triangle Park, NC 27709
William M. Vatavuk
Standards Development Branch, OAQPS
U.S. Environmental Protection Agency
Research Triangle Park, NC 27711
November 1991
>
Contents
8-3
8.1 Introduction
8.1.1 System Efficiencies and Performance 8"*
8-1
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8.2 Process Description 8-4
8.2.1 VOC Condensers 8-7
8.2.2 Refrigeration Unit 8-8
8.2.3 Auxiliary Equipment 8-10
8.3 Design Procedures 8-10
8.3.1 Estimating Condensation Temperature 8-12
8.3.2 VOC Condenser Heat Load 8-13
8.3.3 Condenser Size 8-16
x*
8.3.4 Coolant Flow Rate 8-17
8.3.5 Refrigeration Capacity 8-17
8.3.6 Recovered VOC 8-18
8.3.7 Auxiliary Equipment 8-18
8.3.8 Alternate Design Procedure 8-19
8.4 Estimating Total Capital Investment 8-20
8.4.1 Equipment Costs for Packaged Solvent Vapor Recovery
Systems 8-21
8.4.2 Equipment Costs for Nonpackaged (Custom) Solvent Va-
por Recovery Systems 8-25
8.4.3 Equipment Costs for Gasoline Vapor Recovery Systems 8-26-
8.4.4 Installation Costs 8-28
8.5 Estimating Total Annual Cost 8-30
8.5.1 Direct Annual Costs 8-30
8.5.2 Indirect Annual Costs 8-32
8-2
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8.5.3 Recovery Credit 8-32
8.5.4 Total Annual Cost 8-33
8.6 Example Problem #1 8-33
8.6.1 Required Information for Design 8-33
8.6.2 Equipment Sizing 8-33
8.6.3 Equipment Costs 8-37
8.6.4 Total Annual Cost 8-38
8.7 Example Problem #2 8-40
>-
8.7.1 Required Information for Design 8-40
8.8 Acknowledgments 8-40
Appendix 8A - Properties of Selected Compounds 8-42
Appendix 8B - Documentation for Gasoline Vapor Recovery System
Cost Data 8-45
References 8-49
8.1 Introduction
Condensers in use today may fall in either of two categories: refrigerated or
non-refrigerai£4. Non-refrigerated condensers are widely used as raw mate-
rial and/or product recovery devices in chemical process industries. They_
are frequently used prior to control devices (e.g., incinerators or adsorbers).
ReTngerated condensers are used as air pollution control devices tor treating
emission streams with high VQC concentrations (usually->5.1)00 ppmv) in
applications involving gasoline bulk terminals, storage, etc.
Condensation is a separation technique in which one or more volatile com-
ponents of a va.por mixture are separated from the remaining vapors through
saturation followed by a phase change. The^ phasechange from gas to liquid
8-3
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can be achieved in two ways: (a) the system pressure can be increased at
agjverTternperature, or~Tb)~the temperature may be lowered at a constant
pressure._Jn a two-component system where one of the components is non-
condensible (e.g., air), condensation occurs at dew point (saturation) when
the partial pressure of the volatile compound is equal to its vapor pressure.
The more volatile a compound (i.e., the lower the normal boiling point), the
larger the amount that can remain as vapor at a given temperature; hence
the lower the temperature required for saturation (condensation). Refrigera-
tion is often employed to obtain the low temperatures required for acceptable
removal efficiencies. This chapter is limited to the evaluation of refrigerated
condensation at constant (atmospheric) pressure.
8.1.1 System Efficiencies and Performance
The removal efficiency of a condenser is dependent on the emission stream
characteristics including the nature of the VOC in question (vapor pres-
sure/temperature relationship), VOC concentration, and the type of coolant
used. Any component of any vapor mixture can be condensed if brought
to a low enough temperature and allowed to come to equilibrium. Fig-
ure 8.1 shows the vapor pressure dependence on temperature for selected
compounds.[1] A condenser cannot lower the inlet concentration to levels
below the saturation concentration at the coolant temperature. Removal
efficiencies above 90 percent can be achieved with coolants such as chilled
water, brine solutions, ammonia, or chlorofiuorocarbons, depending on the
VOC composition and concentration level of the emission stream.
8.2 Process Description
Figure 8.2 depicts a typical configuration for a refrigerated surface con-
denser system as an emission control device. The basic equipment required
for a refrigerated condenser system includes a VOC condenser, a refrigera-
tion unit(s), and auxiliary equipment (e.g.. precooler. recovery/storage rank.
pump/blower, and piping).
8-4
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1000.0.
400
100.0 —
40
10.0 _
1
1.0 —
0.1 —
001
441 261 141 55.2 -9 -59
TamoerMure <*F)
•99 -131.7
Figure 8.1: Vapor Pressures of Selected Compounds vs Temperature(i)
-------
Air/VCX;
Vapor In
Precooter Condensale
W«er/VOC
Air/Residual VOC
Discharge
Figure 8.2: Schematic Diagram for a Refrigerated Condenser System
8-6
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Coolant
Inlet
Vapor
Outlet
Vapor
Inlet
m m
/ \
L-
Condensed
VOC
Coolant
Outlet
Figure 8.3: Schematic Diagram of a Shell and Tube Surface Condenser[3]
8.2.1 VOC Condensers
The two most common types of condensers used are surface and contact
condensers.[2] In surface condensers, the coolant does not contact the gas
stream. Most surface condensers in refrigerated systems are the shell and
tube type (see Figure 8.3).[3] Shell and tube condensers circulate the coolant
through tubes. The VOCs condense on the outside of the tubes (i.e., within
the shell). Plate and frame type heat exchangers are also used as condensers
in refrigerated systems. In these condensers, the coolant and the vapor flow
separately over thin plates. In either design, the condensed vapor forms a
film on the cooled surface and drains away to a collection tank for storage.
reuse, or disposal.
In contrast, to surface condensers where the coolant does not contact either
the vapors or the condensate, contact condensers cool the vapor stfeam by
spraying either a liquid at ambient temperature or a chilled liquid directly
into the gas stream.
Spent coolant containing the VOCs from contact condensers usually can-
not be reused directly and can be a waste disposal problem. Additionally.
VOCs in the spent coolant can not be directly recovered without further pro-
cessing. Since the coolant from surface condensers does not contact the vapor
-------
Condenser
High Pressure Side
Expansion
Valve
Low Pressure Side
Evaporator
Compressor
Figure 8.4: Basic Refrigeration Cycle[4]
stream, it is not contaminated and can be recycled in a closed loop. Surface
condensers also allow for direct recovery of VOCs from the gas stream. Tliis
chapter addresses the design and costing of refrigerated surface condenser
systems only.
8.2.2 Refrigeration Unit
The commonly used mechanical vapor compression cycle to produce refrigj
eration consists of four stages: evaporation, compression, condensation, and
expansion (see Figure 8.4).(4l The cycle which is used for single-stage va-
por compression involves two pressures, high and low, to enable a continuous
process to produce a cooling effect. Heat absorbed from the gas stream evap-
orates the liquid coolant (refrigerant). Next, the refrigerant (now in vapor
phase) is compressed to a higher temperature and pressure by the system
compressor. Then, the superheated refrigerant vapor is condensed, reject-
8-8
-------
ing its sensible and latent heat in the condenser. Subsequently, the liquid
refrigerant flows from the condenser through the expansion valve, where its
pressure and temperature are reduced to those in the evaporator, thus com-
pleting the cycle.
The capacity of a refrigeration unit is the rate at which heat is removed,
expressed in tons of refrigeration. One ton of refrigeration is the refrigeration
produced by melting one ton of ice at 32°F in 24 hours. It is the rate of
removing heat equivalent to 12,000 Btu/h or 200 Btu/min. For more details
on refrigeration principles, see References [5] and [6].
For applications requiring low temperatures (below about -30°F), mul-
tistage refrigeration systems are frequently employed.[4] Multistage systems
are designed and marketed in two different types—compound and cascade.
In compound systems, only one refrigerant is used. In a cascade system, two
or more separate refrigeration systems are interconnected in such a manner
that one provides a means of heat rejection for the other. Cascade systems
are desirable for applications requiring temperatures between -50 and -150" F
and allow the use of different refrigerants in each cycle.[4] Theoretically, any
number of cascaded stages are possible, each stage requiring an additional
condenser and an additional stage of compression.
In refrigerated condenser systems, two kinds of refrigerants are used, pri-
mary and secondary. Primary refrigerants are those that undergo a phase
change from liquid to gas after absorbing heat. Examples are ammonia
(R-717), and chlorofluorocarbons such as cidorodiiluoromethane (R-22) or
dichiorodifluoromethane (R-12). Recent concerns about the latter causing
depletion of the ozone layer is prompting development of substitute refriger-
ants.
Secondary refrigerants such as brine solutions act only as heat carriers and
remain in liquid phase. Conventional systems use a closed primary refrigerant
loop that cools the secondary loop through the heat transfer medium in the
evaporator. The secondary heat transfer fluid is then pumped to a VOC
vapor condenser where it is used to cool the air/VOC vapor stream. In some
applications, however, the primary refrigeration fluid is directly used to cool
the vapor stream.
3-9
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8.2.3 Auxiliary Equipment
As shown in Figure 8.2, some applications may require auxiliary equipment
such as precoolers, recovery/storage tanks, pumps/blowers, and piping.
If water vapor is present in the treated gas stream or if the VOC has a
high freezing point (e.g., benzene), ice or frozen hydrocarbons may form on
the condenser tubes or plates. This will reduce the heat transfer efficiency
of the condenser and thereby reduce the removal efficiency. Formation of ice
will also increase the pressure drop across the condenser. In such cases, a
precooler may be needed to condense the moisture prior to the VOC con-
denser. This precooler would bring the temperature of the stream down to
approximately 35 to 40°F, effectively removing the moisture from the gas.
Alternatively, an intermittent heating cycle can be used to melt away ice
build-up. This may be accomplished by circulating ambient temperature
brine through the condenser or by the use of radiant heating coils. If a sys-
tem is not operated continuously, the ice can also be removed by circulating
ambient air.
A VOC recovery tank for temporary storage of condensed VOC prior to
reuse, reprocessing, or transfer to a larger storage tank may be necessary in
some cases. Pumps and blowers are typically used to transfer liquid (e.g.,
coolant or recovered VOC) and gas streams, respectively, within the system.
8.3 Design Procedures
In this section are presented two procedures for designing (sizing) refrigerated
surface condenser systems to remove VOC from air/VOC mixtures. With
the first procedure presented, one calculates the condenser exit temperature
needed to obtain a given VOC recovery efficiency. In the second procedure,
which is the inverse of the first, the exit temperature is given and the recovery
efficiency corresponding to it is calculated.
The first procedure depends on knowledge of the following parameters:
1. Volumetric flow rate of the VOC-containing gas stream:
2. Inlet temperature of the gas stream;
8-10
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Table 8.1: Required Input Data
Data Variable Name Units
Inlet Stream Flow Rate Qin scfm (77°F; 1 atm)
Inlet Stream Temperature -Hn °F
VOC Inlet Volume Fraction 3/voc,tn volume fraction
Required VOC Removal Efficiency ?/ fractional (volume)
Antoine Equation Constants" A,B,C —
Heat of Condensation of the VOC" A#voc Btu/lb-mole
Heat Capacity of the VOCa air Btu/lb-mole-0F
"See Appendix 8A for these properties of selected organic compounds.
3. Concentration and composition of the VOC in the gas stream;
4. Required removal efficiency of the VOC;
5. Moisture content of the emission stream; and
6. Properties of the VOC (assuming the VOC is a pure compound):
• Heat of condensation,
• Heat capacity, and
• Vapor pressure.
\
The design of a refrigerated condenser system requires determination of
the VOC condenser size and the capacity of the refrigeration unit. For a given
VOC removal efficiency, the condensation temperature and the heat load
need to be calculated to determine these parameters. The data necessary to
perform the sizing procedures below as well as the variable names and their
respective units are listed in Table 8.1.
The steps outlined below for estimating condensation tempera! lire and
the heat load apply to a two-component mixture (VOC/'air) in which one
of the two components is considered to be noiicondensible (air). The VOC
component is assumed to consist of a single compound. Also, the emission
stream is assumed to be free of moisture. The calculations are based on the
assumptions of ideal gas and ideal solution to simplify the sizing procedures.
For a more rigorous analysis, see Reference [5].
8-li
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8.3.1 Estimating Condensation Temperature
The temperature necessary to condense the required amount of VOC must be
estimated to determine the heat load. The first step is to determine the VOC
concentration at the outlet of the condenser for a given removal efficiency.
This is calculated by first determining the partial pressure of the VOC at the
outlet assuming that the ideal gas law applies:
Moles VOC in outlet stream fo t\
VOC partial pressure (outlet) = 760— ——~—;T (°-L)
* r ^ Moles inlet stream - Moles VOC removed
However:
Moles VOC in outlet stream = (Moles VOC in inlet stream)(l - T/)(8.2)
Moles VOC in inlet stream = (Moles in inlet stream) yvoc,m (8-3)
Moles VOC remo.ved = (Moles VOC in inlet stream)?/ (8.4)
where
77 = removal efficiency of the condenser system (fractional)
= Moles VOC removed/Moles VOC in inlet
2/voc in — Volume fraction of VOC in inlet stream
After substituting these variables in Equation 8.1, we obtain:
^voc ~ 76u i _ 7. . („} *•-••-'
1 y voc, ml7/;
where
Pvoc — Partial pressure of the VOC in the exit stream (mm Hg).
The condenser is assumed to operate at a constant pressure of one atmosphere
(760 mm Hg).
At the condenser outlet, the VOC in the gas stream is assumed to he at
equilibrium with the VOC condensate. At equilibrium, the partial pressure
of the VOC in the gas stream is equal to its vapor pressure at that temper-
ature assuming the condensate is pure VOC (i.e., vapor pressure = PVoc)-
Therefore, by determining the temperature at which this condition occurs.
the condensation temperature can be specified. This calculation is based on
8-12
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the Antoine equation that defines the relationship between vapor pressure
and temperature for a particular compound:
D
log(vapor pressure) = log Pvoc = A - —— (8.6)
•* rnn + ^
• con
where Tcon is the condensation temperature (°C). Note that Tcon Is in degrees
Centigrade in this equation. In Equation 8.6, A, #, and C are VOC-specific
constants pertaining to temperature expressed in °C and pressure in mm Hg
(see Appendix 8A). Solving for Tcon and converting to degrees Fahrenheit:
The calculation methods for a gas stream containing multiple VOCs are
complex, particularly when there are significant departures from the ideal
behavior of gases and liquids. However, the temperature necessary for con-
densation of a mixture of VOCs can be estimated by the weighted average
of the temperatures necessary to condense each VOC in the gas stream at a
concentration equal to the total VOC concentration.[1]
8.3.2 VOC Condenser Heat Load
Condenser heat load is the amount of heat that must be removed from the
inlet stream to attain the specified removal efficiency. It is determined from
an energy balance, taking into account the enthalpy change due to the tem-
perature change of the VOC, the enthalpy change due to the condensation
of the VOC, and the enthalpy change due to the temperature change of the
air. Enthalpy change due to the presence of moisture in the inlet gas stream
is neglected in the following analysis.
For the purpose of this estimation, it is assumed that the total heat load
on the system is equal to the VOC condenser heat load. Realistically, when
calculating refrigeration capacity requirements for low temperature cooling
units, careful consideration should be given to the process line losses and
heat input of the process pumps. Refrigeration unit capacities are typically
rated in terms of net output and do not reflect any losses through process
pumps or process lines.
8-13
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First, the number of Ib-moles of VOC per hour in the inlet stream must
be calculated by the following expression:
,in = % (yvoc.J 60 (8-8)
where A/Voc in ls ^e molar flow rate of VOC in the inlet stream. The factor
392 is the volume (ft3) occupied by one Ib-mole of inlet gas stream at standard
conditions (77°F and 1 atm). The number of Ib-moles of VOC per hour in
the outlet gas stream is calculated as follows:
Mvoc.ot.fr = Mvoc,in(l ~ l) (8-9)
where Mvoc oui is the molar flow rate of VOC in the exit stream. Finally,
the number of Ib-moles of VOC per hour that are condensed is calculated as
follows:
MVOC.'COH. = MVOC,in ~ ^VOC.out
where Mvoc con is the flow rate of the VOC that is condensed.
The condenser heat load is then calculated by the following equation:
H, j = A// 4- AH + AH (8.11)
load ^"^ COTl it.7i.mn. ' nnncoTl \ /
where
= condenser heat load (Btu/hr)
_ = enthalpy change associated with the condensed
VOC (Btu/hr)
!\Huncon = enthalpy change associated with the uncondensed
VOC (Btu/hr)
!\Hnoncon = enthalpy change associated with the noncondensible
air (Btu/hr).
The change in enthalpy of the condensed VOC is calculated as follows:
= Afvoc.con A//VOC ^ Cp,voc(T>n ~ Tc
where A//voc is the molar heat of condensation and C „ Voc 's ^ie m°lar 'ieat
capacity of the VOC. Ea.ch parameter varies as a function of temperature. In
Equation 8.12. A//"voc and C voc are evaluated at the mean temperature.
-------
The heat of condensation at a specific temperature, T^ (°R), can be cal-
culated from the heat of condensation at a reference temperature, Tj (°R),
using the Watson Equation: [7]
at T2) = (Atfvoc at T.) ~ (8.13)
where Tc (°R) is the VOC critical temperature.
The heat capacity can also be calculated for a specific temperature, T2,
if heat capacity constants (a, 6, c, and
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8.3.3 Condenser Size
Condensers are sized based on the heat load, the logarithmic mean temper-
ature difference between the emission and coolant streams, and the overall
heat transfer coefficient. The overall heat transfer coefficient, [7, can be es-
timated from individual heat transfer coefficients of the gas stream and the
coolant. The overall heat transfer coefficients for tubular heat exchangers
where organic solvent vapors in noncondensible gas are condensed on the
shell side and water/brine is circulated on the tube side typically range from
20 to 60 Btu/hr-ft2-°F according to Perry's Chemical Engineers' Handbook^}.
To simplify the calculations, a single "(/" value may be used to size these
condensers. This approximation is acceptable for purposes of making study
cost estimates.
•V-
Accordingly, an estimate of 20 Btu/hr-ft2-°F can be used to obtain a
conservative estimate of condenser size. The following equation is used to
determine the required heat transfer area:
load
where
A con = condenser surface area (ft2)
U = overall heat transfer coefficient (Btu/hr-ft2-°F)
l — logarithmic mean temperature difference (°F).
The logarithmic mean temperature difference is calculated by the following
equation, which is based on the use of a countercurrent flow condenser:
( *• in ~~ cool, out) ~~ \ con ~ cool, in) / Q , Q \
- (8.18)
i
r.nol. out
i^T1
^ -L con •*• cool, in
where
Tcooiin — coolant inlet temperature ("F)
TCOol out ~ coolant outlet temperature (°F).
The temperature difference ("approach") at the condenser exit can be as-
sumed to be 15° F. In other words, the coolant inlet temperature, Tcooi
-------
temperature rise of the coolant is specified as 25°F. (These two temperatures—
the condenser approach and the coolant temperature rise—reflect good design
practice that, if used, will result in an acceptable condenser size.) Therefore,
the following equations can be applied to determine the coolant inlet and
outlet temperature:
Tcool,in = Tcon-15 (8.19)
Tcooi,oui = Tcool>in + 25 (8.20)
8.3.4 Coolant Flow Rate
X*
The heat removed from the emission stream is transferred to the coolant. By
a simple energy balance, the flow rate of the coolant can be calculated as
follows:
W , = —— (8.21)
Y* cool f (T — T • ~]
p,cool v cool, out cool,in)
where Wcooi is the coolant flow rate (Ib/hr), and CpiCOOi is the coolant spe-
cific heat (Btu/lb-°F). CpiCOol will vary according to the type of coolant used.
For a 50-50 (volume %) mixture of ethylene glycol and water, CpjCOol is ap-
proximately 0.65 Btu/lb-°F. The specific heat of brine (salt water), another
commonly used coolant, is approximately 1.0 Btu/lb-°F.
8.3.5 Refrigeration Capacity
The refrigeration unit is assumed to supply the coolant at the requiced tem-
perature to the condenser. The required refrigeration capacity 'is expressed
in terms of refrigeration tons as follows:
R _ Hload (8.22)
12,000
Again, as explained in Section 8.3.2, Hload does not include any heat losses.
8-17
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8.3.6 Recovered VOC
The mass of VOC recovered in the condenser can be calculated using the
following expression:
x MWVOC (8.23)
where
Wvoc con = mass of VOC recovered (or condensed) (Ib/hr)
= molecular weight of the VOC (Ib/lb-mole).
8.3.7 Auxiliary Equipment
The auxiliary equipment for a refrigerated surface condenser system may
include:
• precooler,
• recovered VOC storage tank,
• pumps/blowers, and
• piping/ductwork.
If water vapor is present in the treated gas stream, a precooler may be
needed to remove moisture to prevent ice from forming in the VOC con-
denser. Sizing of a precooler is influenced by the moisture concentration and
the temperature of the emission stream. As discussed in Section 8.2.3, a pre-
cooler may not be necessary for intermittently operated refrigerated surface
condenser systems where the ice will have time to melt between successive
operating periods.
If a precooler is required, a typical operating temperature is 35 to 40°F.
At this temperature, almost all of the water vapor present will be condensed
without danger of freezing. These condensation temperatures roughly cor-
respond to a removal efficiency range of 70 to 80 percent if the inlet stream
is saturated with water vapor at 77°F. The design procedure outlined in the
8-18
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previous sections for a VOC condenser can be used to size a precooler, based
on the psychrometric chart for the air-water vapor system (see Reference [4]).
Storage/recovery tanks are used to store the condensed VOC when direct
recycling is not a suitable option. The size of these tanks is determined
from the amount of VOC condensate to be collected and the amount of time
necessary before unloading. Sizing of pumps and blowers is based on the
liquid and gas flow rates, respectively, as well as the system pressure drop
between the inlet and outlet. Sizing of the piping and ductwork (length and
diameter) primarily depends upon the stream flow rate, duct/pipe velocity,
available space, and systejn layout.
8.3.8 Alternate Design Procedure
In some applications, it may be desirable to size and cost a refrigerated
condenser system that will use a specific coolant and provide a particular
condensation temperature. The design procedure to be implemented in such
a case would essentially be the same as the one presented in this section
except that instead of calculating the condenser exit temperature needed to
obtain a specified VOC recovery efficiency, the exit temperature is given and
the corresponding recovery efficiency is calculated.
The initial calculation would be to estimate the partial ( = vapor) pressure
of the VOC at the given condenser exit temperature, Tcon, from Equation
8.6. Next, calculate 77 using Equation 8.24, which is Equation 8.5 rearranged:
v760yvoc>HJ-Pvoc ' (8,4)
Finally, substitute the calculated Pvoc m^° ^"s equation to obtain rj. in the
remainder of the calculations to estimate condenser heat load, refrigeration
capacity, coolant flow rate, etc., follow the procedure presented in Sections
8.3.2 through 8.3.7.
8-19
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8.4 Estimating Total Capital Investment
This section presents the procedures and data necessary for estimating capital
costs for refrigerated surface condenser systems in solvent vapor recovery and
gasoline vapor recovery applications. Costs for packaged and nonpackaged
solvent vapor recovery systems are presented in Sections 8.4.1 and 8.4.2, re-
spectively. Costs for packaged gasoline vapor recovery systems are described
in Section 8.4.3. Costs are calculated based on the design/sizing procedures
discussed in Section 8.3.
Total capital investment, TCI, includes equipment cost, EC, for the entire
refrigerated condenser unit, auxiliary equipment costs, taxes, freight charges,
instrumentation, and direct and indirect installation costs. All costs in this
chapter are presented 3rd quarter 1990 dollars.
For these control systems, the total capital investment is a battery limit
cost estimate and does not include the provisions for bringing utilities, ser-
vices, or roads to the site; the backup facilities; the land; the working capital;
the research and development required; or the process piping and instrumen-
tation interconnections that may be required in the process generating the
waste gas. These costs are based on new plant installations; no retrofit cost
considerations are included. The retrofit cost factors 'are so site specific that
no attempt has been made to provide them.
The expected accuracy of the cost estimates presented in this chapter is
±30 percent (i.e., "study" estimates). It must be kept in mind that even for
a given application, design and manufacturing procedures vary from vendor
to vendor, so costs may vary. ,
In the next two sections, equipment costs are presented for packaged nnd
nonpackaged (custom) solvent vapor recovery systems, respectively. With
the packaged systems, the equipment cost is factored from the refrigeration
unit cost: with the custom systems, the equipment cost is determined as the
sum of the costs of the individual system components. Finally, equipment
costs for packaged gasoline vapor recovery systems are given in Section 8.4.3.
8-20
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8.4.1 Equipment Costs for Packaged Solvent Vapor
Recovery Systems
Vendors were asked to provide refrigerated unit cost estimates for a wide
range of applications. The equations shown below for refrigeration unit
equipment costs, ECr, are multivariable regressions of data provided by two
vendors and are only valid for the ranges listed in Table 8.2.[8, 9] In this table,
the capacity range of refrigeration units for which cost data were available
are shown as a function of temperature.
Single Stage Refrigeration Units (less than 10 tons)
ECr = exp (9.83 - 0.014Tcon + 0.340In R) (8.25)
X-
Single Stage Refrigeration Units (greater than or equal to 10 tons)
ECr = exp (9.26 - 0.007Tcon + 0.627In R) (8.26)
Multistage Refrigeration Units
ECr = exp (9.73 - 0.012Tcon + 0.584 In R) (8.27)
NOTE: exp(a) = ea ss 2.718"
Equations 8.25 and 8.26 provide costs for refrigeration units based on
single stage designs, while Equation 8.27 gives costs for multistage units.
Equation 8.27 covers both types of multistage units, "cascade" and "com-
pound". Data provided by a vendor show that the costs of cascade and
compound units compare well, generally differing by less than 30%.[8] Thus,
only one cost equation is provided. Equation 8.25 applies to single stage
refrigeration units smaller than 10 tons and Equation 8.26 applies to single
stage refrigeration units as large or larger than 10 tons. Single stage units
typically achieve temperatures between 40 and -20° F. although there are
units that are capable of achieving -60°F in a single stage.[8. 10j Multistage
units are capable of lower temperature operation between -10 and -100° F.
Single stage refrigeration unit costs are depicted graphically for selected
temperatures in Figure 8.5. Figure 8.6 shows the equipment cost curves for
multistage refrigeration units.
8-21
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Table 8.2: Applicability Ranges for the Refrigeration Unit Cost Equations
(Equations 8.25 to 8.27)
Temperature
T (TV
-1 con \ *• I
40
30
20
10
0 to -5
-10
-20 to -25
-30
-40
-45 to -50
-55 to -60
-70
-75 to -80
-90
-100
Minimum Size Available
R (tons)
Single Stage
0.85
0.63
0.71
0.44 " '
0.32
0.21
0.13
NA
NA
NA
NA
NA
NA
NA
NA
Multistage
NA6
NA
NA
NA
NA
3.50
2.92
2.42
1.92
1.58
1.25
1.33
1.08
0.83
0.67
Maximum Size Available
R (tons)
Single Stage
174
170
£80
200
133
6.6
200
NA
NA
100"
100C
NA
NA
Multistage
NA
NA
NA
NA
NA
81
68
85
68
55
100
42
150
NA 28
NA 22
"For condensation temperatures that lie between the levels shown,
round off to the nearest level (e.g., if Tcon = 16°F, use 20°F) to de-
termine minimum and maximum available size.
6NA = System not available based on vendor data collected in this
study.
'Only one data point available.
8-22
-------
200,000
160,000
<§
- 120,000
«
o
8
O 80,000
«i
•
_
3
cr
IU
40.000
or..
20-F
40°F
-20T .'
20
4O
Capacity (tons)
60
80
100
Figure 8.5: Refrigeration Unit Equipment Cost (Single Stage)[8, 9]
NOTE: The discontinuities in the curves at the 10 ion capacity are a result of the
two regression equations used. Equation 8.25 is used for capacities of less than 10
tons; Equation 8.26 is used for capacities greater rhan or equal to 10 tons.
8-23
-------
700,000
600,000
500,000
400,000
300,000
g. 200,000
100,000
O
8
o>
1
-eo-F
-60-F
-100-F.-
-40-F
..---:' -30-F
-20-F
20
40 60
Capacity (tons)
30
100
Figure 8.6: Refrigeration Unit Equipment Cost (Multistage)[8, Ol
8-24
-------
These costs are for outdoor models that are skid-mounted on steel beams
and consist of the following components: walk-in weatherproof enclosure,
air-cooled low temperature refrigeration machinery with dual pump design,
storage reservoir, control panel and instrumentation, vapor condenser, and
necessary piping. All refrigeration units have two pumps: a system pump
and a bypass pump to short-circuit the vapor condenser during no-load con-
ditions. Costs for heat transfer fluids (brine) are not included.
The equipment cost of packaged solvent vapor recovery systems (EC/«)
is estimated to be 25 percent greater than the cost of the refrigeration unit
alone [9]. The additional cost includes the VOC condenser, recovery tank,
the necessary connections, piping, and additional instrumentation. Thus,
EC/> = 1.25ECr (8.28)
Purchased equipment cost, PEC, includes the packaged equipment cost,
and factors for .sales taxes (0.03) and freight (0.05). Instrumentation
and controls are included with the packaged units. Thus,
PECp = ECr(l + 0.03 + .05) = 1.08ECF (8.29)
8.4.2 Equipment Costs for Nonpackaged (Custom)
Solvent Vapor Recovery Systems
To develop cost estimates for nonpackaged or custom refrigerated systems,
information was solicited from vendors on costs of refrigeration units, VOC
condensers, and VOC storage/recovery tanks [9, 11, 12]. Quotes from the
vendors were used to develop the estimated costs for each type of equipment.
Only one set of vendor data was available for each type of equipment.
Equations 8.25, 8.26, and 8.27 shown above are applicable for estimating
the costs for the refrigeration units.
Equation 8.30 shows the equation developed for the VOC condenser cost
estimatesflll:
34.4con^,775 (8-30)
This equation is valid for the range of 38 to 800 iV and represents costs for
shell and tube type heat exchangers with 304 stainless steel tubes.
8-25
-------
The following equation represents the storage/recovery tank cost data
obtained from one vendor[12]:
£0^ = 2.72^+1,960 (8.31)
These costs are applicable for the range of 50 to 5,000 gallons and pertain to
316 stainless steel vertical tanks.
Costing procedures for a precooler (ECprc) that includes a separate con-
denser/refrigeration unit and a recovery tank are similar to that for a custom
refrigerated condenser system. Hence, Equations 8.25 through 8.31 would be
applicable, with the exception of-Equation 8.27, which represents multistage
systems. Multistage systems operate at much lower temperatures than that
required by a precooler.
•\-
Costs for auxiliary equipment such as ductwork, piping, fans, or pumps
are designated as ECawz. These items should be costed separately using
methods described elsewhere in this Manual.
The total equipment cost for custom systems, ECc, is then expressed as:
ECC = ECr + ECcon + ECiank + ECpre + ECaui (8.32)
The purchased equipment cost including ECc and factors for sales taxes
(0.03), freight (0.05), and instrumentation and controls (0.10) is given below:
PECC = ECC(1 + 0.03 + 0.05 + 0.10) = 1.18ECr (8.33)
8.4.3 Equipment Costs for Gasoline Vapor Recovery
Systems
Separate quotes were obtained for packaged gasoline vapor recovery systems
because these systems are specially designed for controlling gasoline vapor
emissions from such sources as storage tanks, gasoline bulk terminals, and
marine vessel loading and unloading operations. Systems that control marine
vessel gasoline loading and unloading operations also must meet U.S. Coast
Guard safety requirements.
Quotes obtained from one vendor were used to develop equipment cost
estimates for these packaged systems (see Figure 8.7). The cost equation
8-26
-------
IB
5
6
•o
O
o>
£.
cr
HI
800,000
600,000
400,000
200,000
Capacity (gal/inln)
Capacity (lono)
2,000
20
_L
_L
4,000
6,000
8,000
40
60
80
100
10,000
120
Figure 8.7: Gasoline Vapor Recovery System Equipment Cost[9]
-------
shown below is a least squares regression of these cost data and is valid for
the range 20 to 140 tons.[9]
ECP = 4,910fl + 212,000 (8.34)
The vendor data in process flow capacity (gal/min) vs cost ($) were trans-
formed into Equation 8.34 after applying the design procedures in Section 8.3.
Details of the data transformation are given in Appendix 8B.
The cost estimates apply to skid-mounted refrigerated VOC condenser
systems for hydrocarbon vapor recovery primarily at gasoline loading/storage
facilities. The systems are intermittently operated at -80 to -120° F allow-
ing 30 to 60 minutes per day for defrosting by circulation of warm brine.
Multistage systems are employed to achieve these lower temperatures. The
achievable VOC removal efficiencies for these systems are in the range of 90
to 95 percent.
The packaged systems include the refrigeration unit with the necessary
pumps, compressors, condensers/evaporators, coolant reservoirs, the VOC
condenser unit and VOC recovery tank, precooler, instrumentation and con-
trols, and piping. Costs for heat transfer fluids (brines) are not included. The
purchased equipment cost for these systems includes sales tax and freight and
is calculated using Equation 8.29.
8.4.4 Installation Costs
The total capital investment. TCI, for packaged systems is obtained by
multiplying the purchased equipment cost, PECp, by the total installation
factor:f!3]
TCI = l.lSPECp (8.35)
For nonpackaged (custom) systems, the total installation factor is 1.74:
TCI - i.74PECr (8.36)
An itemization of the total Installation factor for nonpackaged systems is
shown in Table 3.3. Depending on the site conditions, the installation costs
for a given system could deviate significantly from costs generated by these
average factors. Guidelines are available for adjusting these average installa-
tion factors.[14]
8-28
-------
Table 8.3: Capital Cost Factors for Nonpackaged (Custom) Refrigerated
Condenser Systems
Cost Item
Factor
Direct Costs
Purchased equipment costs
Refrigerated condenser system, EC
Instrumentation
Sales taxes
Freight
Purchased equipment cost, PEC
Direct installation costs
Foundations & supports
Handling & erection
Electrical
Piping
Insulation
Painting
Direct installation costs
Site preparation
Buildings
Total Direct Costs, DC
Indirect Costs (installation)
Engineering
Construction and field expenses
Contractor fees
Start-up
Performance test
Contingencies
Total Indirect Costs, 1C
As estimated, A
0.10 A
0.03 A
0.05 A
B~^~1.18 A"
0.08 B
0.14 B
0.08 B
0.02 B
0.10 B
0.01 D
(T43~B
As required, SP
As required, Bldg.
1.43 B + SP + Bldg.
0.10 B
0.05 B
0.10 B
0.02 B
0.01 B
0.03 B
0.31 B
Total Capital Investment = DC + 1C
1.74 B -I- SP + Bldg.'' !
"Purchased equipment cost factor for packaged systems is 1.08 with
instrumentation included.
6For packaged systems, total capital investment = 1.15PEC/-.
8-29
-------
8.5 Estimating Total Annual Cost
The total annual cost (TAG) is the sum of the direct and indirect annual
costs. The bases used in calculating annual cost factors are given in Table 8.4.
8.5.1 Direct Annual Costs
Direct annual costs, DC, include-labor (operating and supervisory), mainte-
nance (labor and materials), and electricity.
Operating labor is estimated at 1/2-hour per 8-hour shift. The supervi-
sory labor cost is estimated at 15% of the operating labor cost. Maintenance
labor is estimated at 1/2-hour per 8-hour shift. Maintenance materials costs
are assumed to equal maintenance labor costs.
Utility costs for refrigerated condenser systems include electricity require-
ments for the refrigeration unit and any pumps/blowers. The power required
by the pumps/blowers is negligible when compared with the refrigeration
unit power requirements. Electricity requirements for refrigerated condenser
systems are summarized below:
Electricity (E, kW/ton) Temperature (°F)
1.3
2.2
4.7
5.0
11.7
40
20
-20
-50
-100
These estimates were developed from product literature obtained from one
vendor. [9] The electricity cost, C, ., can then be calculated from the following
expression:
where
$, = system operating hours (hr/yr)
Pr — electricity cost ($/kWh).
8-30
-------
Table 8.4: Suggested Annual Cost Factors for Refrigerated Condenser Systems
Cost Item
Direct Annual Costs, DC
Operating Labor
Operator
Supervisor
Operating materials
Maintenance
Labor
Material
Electricity[9]
at 40°F
at 20°F
at -20°F
at -50°F
at -100°F
Indirect Annual Costs, 1C
Overhead
Administrative charges
Property tax
Insurance
Capital recovery"
Recovery Credits, RC
Recovered VOC
Total Annual Cost
Factor
1/2 hour per shift
15% of operator
1/2 hour per shift
100% of maintenance labor
1.3 kW/ton
2.2 kW/ton
4.7 kW/ton
5.0 kW/ton
11.7 kW/'-m
60% of total labor and
maintenance material costs
2% of Total Capital Investment
1% of Total Capital Investment
1% of Total Capital Investment
0.1315 x Total Capital Investment
Quantity recovered x operating hours
DC 4- 1C- RC
"Assuming a 15 year life at 10%.[13j See Chapter 2.
8-31
-------
The factor 0.85 accounts for the mechanical efficiency of the compressor.[1]
8.5.2 Indirect Annual Costs
Indirect annual costs, 1C, are calculated as the sum of capital recovery costs
plus general and administrative (G&A), overhead, property tax, and insur-
ance costs. Overhead is assumed to be equal to 60 percent of the sum of
operating, supervisory, and maintenance labor, and maintenance materials.
Overhead cost is discussed in Chapter 2 of this Manual.
The system capital recovery cost, CRC, is based on an estimated 15-
year equipment life.[13] (See Chapter 2 of the Manual for a discussion of the
capital recovery cost.) For a'15-year life and an interest rate of 10 percent,
the capital recovery factor is 0.1315. The system capital recovery cost is then
estimated by:
CRC = 0.1315 TCI (8.38)
GfeA costs, property tax, and insurance are factored from total capital
investment, typically at 2 percent, 1 percent, a.nd 1 percent, respectively.
8.5.3 Recovery Credit
If the condensed VOC can be directly reused or sold without further treat-
ment, then the credit from this operation is incorporated in the total annual
cost estimates. The following equation can be used to estimate the VOC
recovery credit, RC:
l.Pvoc (8-39)
where
Pvoc ~ resale value of recovered VOC (S/lb)
^ voc con ~ quantity of VOC recovered (Ib/hr).
8-32
-------
8.5.4 Total Annual Cost
The total annual cost (TAG) is calculated as the sum of the direct and indirect
annual costs, minus the recovery credit:
TAG = DC + 1C - RC (8.40)
8.6 Example Problem
The example problem described in this section shows how to apply the re-
frigerated condenser system sizing and costing procedures to the control of a
vent stream consisting of acetone, air, and a"negligible amount of moisture.
This example problem assumes a required removal efficiency and calculates
the temperature needed to achieve this level of control.
8.6.1 Required Information for Design
The first step in the design procedure is to specify the gas stream to be
processed. Gas stream parameters to be used in this example problem are
listed in Table 8.5. The values for the Antoine equation constants, heat of
condensation, and heat capacity of acetone are obtained from Appendix 8A.
Specific heat of the coolant is obtained from Perry's Chemical Engineers'
Handbook^}.
8.6.2 Equipment Sizing
*
The first step in refrigerated condenser sizing is determining the partial pres-
sure of the VOC at the outlet of the condenser for a given removal efficiency.
Given the stream flow rate, inlet VOC concentration, and the required re-
moval efficiency, the partial pressure of the VOC at the outlet can be calcu-
lated using Equation 8.5.
0.375(1 -0.90) _
1 -0.375(0.90) ~
8-33
-------
Table 8.5: Example Problem Data
Vent Stream Parameters
Inlet Stream Flow Rate 100 scfrn"
Inlet Stream Temperature " 86° F
VOC to be Condensed Acetone
VOC Inlet Volume Fraction 0.375
Required VOC Removal Efficiency .90
Antoine Equation Constants for Acetone:
A 7.117
B 1210.595
C 229.664
Heat of Condensation of Acetone'' 12,510 Btu/lb-mole
Heat Capacity of Acetonec 17.90 Btu/lb-mole-0F
Specific Heat of Coolantc (ethylene glycol) 0.65 Btu/lb-°F
Heat Capacity of Airc 6.95 Btu/lb-mole-°F
Annual Cost Data Assumed
Operating Labor $15.64/hr
Maintenance Labor $17.21/hr
Electricity $0.0461/kWh
Acetone Resale Value $0.10/lb
"Standard conditions: 77°F and 1 atmosphere.
bEvaluated at the acetone boiling point (134°F).
''These properties were evaluated at 77°F.
8-34
-------
Next, the temperature necessary to condense the required amount of VOC
must be determined using Equation 8.7:
I "iQ1! \
- 229.664 I 1.8 + 32 = 16°F
u,
43
The next step is to estimate the VOC condenser heat load. Calculate:
(1) the VOC flow rate for the inlet/outlet emission streams, (2) the flow rate
of the condensed VOC, and (3) the condenser heat balance. The flow rate of
VOC in the inlet stream is calculated from Equation 8.8.
^vocin = (0-375) 60 = 5.74 Ib-moles/hr
'VOC,tn
The flow rate of VOC in the outlet stream is calculated using Equation 8.9
as follows:
Mvoc.out = S-74^ ~ °-9°) = °-574 lb-moles/hr
Finally, the flow rate of condensed VOC is calculated with Equation 8.10:
AIvoc,con = 5'74 ~ °'574 = 5'166 lb-moles/hr
Next, the condenser heat balance is conducted. As indicated in Table 8.5,
the acetone heat of condensation is evaluated at its boiling point, 134°F.
However, it is assumed (for simplicity) that all of the acetone condenses
at the condensation temperature, Tcon = 16°F. To estimate the heat of
condensation at 16°F, use the Watson equation (Equation 8.13) with the
following inputs: Tr = 918°R(Appendix 8A); T, = 134 -f 460 - 594°R;
T2 = 16 + 460 = 476°R. Upon substitution, we obtain:
/I- 476/918 V'"''8
(Atfvoc at 16«F) = 12,510^_594'/918j
= 1-1,080 Btu/'lb-moie
As Table 8.5 shows, the heat capacities of acetone and air and the specific
heat of the coolant were all evaluated at 77° F. This temperature is fairly
close to the condenser mean operating temperature, i.e., (86 -r 16J/2 = 51°F.
Consequently, using the 77°F values would not add significant additional
error to the heat load calculation.
8-35
-------
The change in enthalpy of the condensed VOC is calculated using Equa-
tion 8.12:
&Hcon = 5.166 [14,080 + 17.90(86 - 16)] = 79,210 Btu/hr
The enthalpy change associated with the uncondensed VOC is calculated
from Equation 8.15:
&Huncon = (0.574)(17.90)(86 - 16) = 719 Btu/hr
Finally, the enthalpy change of the noncondensible air is estimated from
Equation 8.16:
~ 5'74
The condenser heat load is then calculated by substituting AjFfcon, A/funcon,
and &Hnoncon in Equation 8.11:
Blood = 79> 21° + 719 + 4' 654 = 84'583 Btu/hr
The next step is estimation of the VOC condenser size. The logarithmic
mean temperature difference is calculated using Equation 8.18. In this calcu-
lation, Tcooliin = 16- 15 = 1°F and Tcooli0ut = 1 + 25 = 26°F from Equations
8.19 and 8.20, respectively:
(86_26)-(16-_L) =32>5<,F
16 - 1
The condenser surface area can then be calculated using Equation 8.11
84,583
In this equation, a. conservative value of 20 Btu/hr-ft'-'-°F is used as the overall
heat transfer coefficient.
The coolant flow rate can be calculated using Equation 8.21.
\V = 84,583 =
cooi 0.65(26 - 1)
8-36
-------
The refrigeration capacity can be estimated from Equation 8.22 as follows:
Finally, the quantity of recovered VOC can be estimated using Equation 8.23:
WVOC.OTO = 5-166 x 58'08 = 30° lb/hr
where the molecular weight of acetone is obtained from Appendix 8A.
Note that in this example case, the partial pressure of acetone at the con-
denser exit is relatively high (43 mm Hg). In applications where much lower
outlet concentrations are desired, a second control device (e.g., incinerator,
adsorber) to operate in series with the condenser may need to be considered.
>•«
8.6.3 Equipment Costs
Once the system sizing parameters have been determined, the equipment
costs can be calculated. For the purpose of this example, a custom refriger-
ated condenser system, including a refrigeration unit, a VOC condenser, and
a recovery tank will be costed.
From Table 8.2, a single stage refrigeration unit appears to be suitable for
the example problem with an estimated condensation temperature of 16° F
and capacity of 7.05 tons. Hence Equation 8.25, which is applicable to units
less than 10 tons, is selected for estimating costs. Application of this equation
results in the following value for the refrigeration unit cost:
ECr = exp[9.83 -0.014(16) + 0.340 ln(7.05)] = 328,855
VOC condenser cost is computed using Equation 8.30 as follows:
ECcon, = 34(130) + 3, 775 = $8, 195
Recovery tank cost can he calculated from Equation 8.31. For lliis case,
Wvoc con — 300 lb/hr. which is equivalent to 45.5 gal/hr (density of acetone
is about 6.6 Ib/gal). Assuming an 8-hour daily operation, the interim storage
capacity requirement would be 364 gallons. Application of Equation 8.31
leads to the following:
ECian/fe = --"2(364) f 1,960 = $2,950
8-37
-------
Assuming there are no additional costs due to precooler or other auxiliary
equipment, the total equipment cost is calculated from Equation 8.32:
ECc = 28,855 + 8,195 + 2,950 + 0 + 0 = $40,000
The purchased equipment cost including instrumentation, controls, taxes,
and freight is estimated using Equation 8.33:
PECc = 1.18(40,000) = $47,200
The total capital investment is calculated using Equation 8.36:
TCI = 1.74(47,200) = $82,128 w $82,100
8.6.4 Total Annual Cost
Table 8.6 summarizes the estimated annual costs for the example problem.
The cost calculations are shown in the table.
Direct annual costs for refrigerated systems include labor, materials, and
utilities. Labor costs are based on 8-hr/day, 5-day/week operation. Supervi-
sory labor is computed at 15 percent of operating labor, and operating and
maintenance labor are each based on 1/2 hr per 8-hr shift. The electricity
cost is based on a requirement of 2.2 kW/ton (see page 8-30), because the
condensation temperature (16°F) is close to the 20°F temperature given for
this value.
Indirect annual costs include overhead, capital recovery, administrative
charges, property tax. and insurance.
Total annual cost is estimated using Equation 8.40. For this example
case, application of refrigerated condensation as a control measure results
in an annual savings of approximately $36,000. As Table 8.6 shows, the
acetone recovery credit is over twice the direct and indirect costs combined.
Clearly, this credit has more influence on the total annual cost than any other
component. Although the credit depends on three parameters—the acetone
recovery rate, the annual operating hours, and the acetone salvage value
($0.10/lb)—the last parameter is often the most difficult to estimate. This
is mainly because the salvage value varies according to the facility location
as well as the current state of the chemical market.
8-38
-------
Table 8.6: Annual Cost for Refrigerated Condenser System
Example Problem
Cost Item
Direct Annual Costs, DC
Operating Labor
Operator
Supervisor
Operating materials
Maintenance
Labor
Material
Utilities
Electricity
Total DC
Indirect Annual Costs, 1C
Overhead
Administrative charges
Property tax
Insurance
Capital recovery"
Total 1C
Recovery Credits, RC
Recovered Acetone
Calculations
0.5 h v shift v 2,080 h $15.64
1 'A, X ~lf"tr~ •* — yr C
15% of operator = 0.15 x 2,030
0.5 h v shift v 2,080 h. $17.21
"slufT 8 h yr R
100% of maintenance labor
7.05 tons ,, 2.2 kW 2,080 h $0.0461
0.85 x ton x yr x kWh
60% of total labor and maintenance material:
= 0.6(2,030 + 305 + 2,240 + 2,240)
2% of Total Capital Investment = 0.02(882,100)
1% of Total Capital Investment = 0.01(882,100)
1% of Total Capital Investment = 0.01(882,100)
0.1315 x $82,100
300 Ib x 2,080 h x 80.10
Total Annual Cost (rounded)
Cost
$2,030
300
—
2,240
2,240
1,750
$8,560
4,090
1,640
820
820
lOfSOO
SL8,170
(862,400)
($36,000)
(Savings)
"The capital recovery cost factor, CRF, is a function of the refrigerated condenser equip-
ment life and the opportunity cosi of the capital (i.e.. interest rate). For example, for a
15 year equipment life and a 10% interest rate, CRF = 0.1315.
8-39
-------
8.7 Example Problem #2
In this example problem, the alternate design procedure described in Sec-
tion 8.3.8 is illustrated. The temperature of condensation is given, and the
resultant removal efficiency is calculated. The example stream inlet param-
eters are identical to Example Problem #1 with the exception that removal
efficiency is not specified and the required temperature of condensation is
assumed to be 16°F.
8.7.1 Required Information for Design
The first step is to calculate the partial pressure of the VOC at the specified
temperature (16°F) using Equation 8.6 to solve for ?Voc:
B
= A ~ - -
con
con
Remember to convert Tcon to degrees Centigrade, i.e., 16°F = -8.9°C.
Substituting the values for the Antoine equation constants for acetone as
listed in Table 8.5:
1210.595
logPvoc =
-8.9 + 229.664
pvoc = 43 mm
Using Equation 8.24. the removal efficiency is:
[760(0.375)j- 43 '
n~ 0.375(760-43)
The remainder of the calculations in this problem are identical to those in
Example Problem #1.
8.8 Acknowledgments
The a.uthors gratefully acknowledge the following companies lor contributing
data to this chapter:
8-40
-------
-------
• Edwards Engineering (Pompton Plains, NJ)
• Piedmont Engineering (Charlotte, NC)
• Universal Industrial Refrigeration (Gonzales, LA)
• ITT Standard (Atlanta, GA)
• XChanger (Hopkins, MN)
• Buffalo Tank Co. (Jacksonville, FL)
8-41
-------
-------
Appendix 8A
Properties of Selected
Compounds
8-42
-------
-------
Table 8.7: Properties of Selected Compounds
Compound
Acetone
Acetylene
Acrylonitrile
Aniline
Benzene
Benzonitrile
Butane
Chloroethane
Chloroform
Chloromethane
Cyclobutane
Cyclohexane
Cyclopentane
Cyclopropane
Diethyl ether
Dimethylamine
Ethylbenzene
Ethylene oxide
Heptane
Hexane
Methanol
Octane
Pentane
Toluene
o - Xylene
TO- Xylene
Critical
Temp."
918
555
—
1259
1012
1259
766
829
966
750
—
997 -'
921
716
840
788
1111
845
973
914
923
1024
846
1065
1135
1111
p- Xylene 1109
Boiling
Point
134
-119
171
364
176
.376
31
54
143
-12
55
177
121
-27
94
44
277
51
209
156
148
258
97
231
292
282
281
Molecular
Weight
(Ib/lb-mole)
58.08
26.02
53.06
93.13
78.11
103.12
58.12
64.52
119.39
50.49
56.10
84.16
70.13
42.08
74.12
45.09
106.17
44.05
100.12
86.18
32.04
114.23
72.15
92.14
106.17
106.17
106,17
Heat of
Condensation6
(Btu/lb-mole)
12,510
7,290
14,040
19,160
13,230
19,800
9,630
10,610
12,740
9,260
10,410
12,890
11,740
8,630
11,480
11,390
15,300
10,980
13,640
12,410
14,830
14.810
11,090
14,270
15,840
15,640
15.480
Heat
Capacity11
/ Btu "\
Ub-mole-°F/
30.22
17.90
10.50
15.24
45.90
25.91
19.52
26.07
23.29
14.97
15.63
9.74
17.26
37.4
25.40
30.80
19.84
13.37
40.8
26.89
16.50
30.69
11.54
53.76
39.67
45.2
34.20
19.40
10.49
45.14
28.73
37.58
24.77
44.9 ,
31.85
43.8
30.49
30.32
State
Liquid
Gas
Gas
Gas
Liquid
Gas
Liquid
Gas
Gas
Gas
Gas
Gas
Gas
Liquid
Gas
Liquid
Gas
Gas
Liquid
Gas
Gas
Gas
Gas
Liquid
Gas
Liquid
'Gas
Liquid
Gas
Gas
Gas
Liquid
Gas
Liquid
Gas
Liquid
Gas
Gas
"Reprinted with permission from Lange's Handbook of Chemistry (12"'
fcReprinted with permission from Lange's Handbook of Chemistry (12"'
edition), Table 9-7 [151
edition), Table 9-4.[15J
(Measured at boiling point.)
' Reprinted with permission from Lange '5 Handbook of Chemistry (12"' edition), Table 9-2 ! !5
(Measured at 77°F.)
8-43
-------
Table 8.8: Antoine Equation Constants for Selected Compounds"
Compound
Acetone
Acetylene
Acrylonitrile
Aniline
Benzene
Benzonitrile
Butane
Chloroethane
Chloroethylene
Chloroform
Chloromethane
Cyanic acid
Cyclobutane
Cyclohexane
Cyclopentane
Cyclopropane
Diethyl ether
Diethylamine
Dimethylamine
Dioxane -1,4
Ethyl benzene
Ethylene oxide
Heptane
Hexane
Methanol
Octane
Pentane
Toluene
Vinyl acetate
Antoine
A
7.117
7.100
7.039
7.320
6.905 .
6.746
6.809
6.986
6.891
6.493
7.093
7.569
6.916
6.841
6.887
6.888
6.920
5.801
7.082
7.432
6.975
7.128
6.897
6.876
7.897
6.919
6.853
6.955
7.210
o - Xviene 6.999
m- Xylene
p - Xylene
7.009
6.991
Equation Constants
B
1210.595
711.0
1232.53
1731.515
1211.033
1436.72
935.86
1030.01
905.01
929.44
948.58
1251.86
1054.54
1201.53
1124.16
856.01
1064.07
583.30
960.242
1554.68
1424.255
1054.54
1264.90
1171.17
1474.08
1351.99
1064.84
1344.8
1296.13
1474.679
1462.266
1453.430
C
229.664
253.4
222.47
206.049
220.790
181.0
238.73
238.61
239.48
196.03
249.34
243.79
241.37
222.65
231.36
246.50
228.80
144.1
221.67
240.34
213.21
237.76
216.54
224.41
229.13
209.15
233.01
Valid
Temperature
Range (°F)
Liquid
-116 to -98
-4 to 248
216 to 365
46 to 217
Liquid
-107 to 66
-69 to 54
-85 to 9
-31 to 142
-103 to 23
-105 to 21
-76 to 54
68 to 178
-40 to 162
-130 to -26
-78 to 68
88 to 142
-98 to 44
68 to 221
79 to 327
-56 to 54
28 to 255
-13 to 198
7 to 149
66 to 306
-58 to 136
219.48 43 to 279
226.66 72 to 162
213.69 90 to 342
215.11 82 to 331
215.31
81 to 331
"Reprinted with permission
(12"' edition), Table 10-8.[15]
from Lanae 's Handbook of Chemistry
8-44
-------
Appendix 8B
Documentation for Gasoline
Vapor Recovery System Cost
Data
As mentioned in Section 8.4.3, vendor cost data were obtained that related
the equipment cost ($) of packaged gasoline vapor recovery systems to the
process flow capacity (gal/min). These data needed to be transformed, in
order to develop Equation 8.34, which relates equipment cost (S) to system
refrigeration capacity (#, tons), as follows:
ECP = 4,910^ + 212,000
To make this transformation, we needed to develop an expression relating
flow capacity to refrigeration capacity. The first step was to determine -the
inlet partial pressure (Pvoc m) of the VOC—gasoline, in this case. As was
done in Section 8.3.1, we assumed that the VOC vapor was saturated and.
thus, in equilibrium with the VOC liquid. This, in turn, meant that we could
equate the partial pressure to the vapor pressure. The "model" gasoline had
a Reid vapor pressure of 10 and a molecular weight of 66 Ib/lb-mole. as
shown in Section 4.3 of Compilation of Air Pollutant Emission Factors (EPA
publication AP-42, Fourth Edition, September 1985). For this gasoline, die
following Antoine equation constants were used:
A = 12.5733
B = 6386.1
C = 613
8-45
-------
These constants were obtained by extrapolating available vapor pressure vs.
temperature data for gasoline found in Section 4.3 of AP-42. Upon substi-
tuting these constants and an assumed inlet temperature of 77°F (25°C) into
the Antoine equation and solving for the inlet partial pressure (PVoc in) we
obtain:
D
= 12.5733-
25 + 613
pvoc,in •= 366 mm HS-
If the system operates at atmospheric pressure (760 mm Hg), this partial
pressure would correspond to a VOC volume fraction in the inlet stream of:
•v
366 mm
The outlet partial pressure (Pvoc oui} and volume fraction are calculated in
a similar way. The condensation (outlet) temperature used in these calcu-
lations is -80°F (-62°C), the typical operating temperature for the gasoline
vapor recovery units for which the vendor supplied costs.
log Pvnr m,y = 12.5733 -- 86'1
6 voco«<
Pvoc.oui - 9-62 mm HS-
This corresponds to a volume fraction in the outlet stream (j/Voc out) °^:
9.62 mm
»— < = Wl^T = °'0127
Substitution of -PVoc out an<^ ^voc in m*° Equation 8.24 yields the condenser
removal efficiency (77):
_ (760 x 0.482) - 9.62 _
11 ~~ 0.482(760 - 9.62) ~
The next step is determining the inlet and outlet VOC hourly molar flow rates
f/]/voc in and .'l/voc out i. respectively). As Equation 3.8 shows. -l/vor in is
a function of yvoc m and the total inlet volumetric flow rate, Q^n (scfm).
8-46
-------
Now, because the gasoline vapor flow rates are typically expressed in gal-
lons/minute, we have to convert them to scfm. This is done as follows:
Qin = Q9 (gal/min) x -^— = 0.134Q, scfm
Substituting these variables into Equation 8.8, we obtain:
(0-482)60 = 0.00989out = 0.00989g9(l - 0.986) = (1.38 x W~l)Qg Ib-moles/hr
And according to Equation 8.10, the amount of gasoline vapor condensed
(MVOC)Con) is the difference between MVOCjjn and MVOCi0ut :
MVOC,con = 0.00975gg Ib-moles/hr
The final step is to calculate the condenser heat load. This load is a function
of the inlet, outlet, and condensate molar flowrates, the inlet and conden-
sation temperatures, the heat capacities of the VOC and air, and the VOC
heat of condensation. The VOC heat capacity and heat of condensation
data used are based on pentane and butane chemical properties, the largest
components of gasoline, and were obtained from CHRIS Hazardous Chemical
Data (U.S. Coast Guard, U.S. Department of Transportation, June 1985).
Heat capacities (Btu/lb-mole-°F):
CP,voc = 26'6
Cp,«r= 6'95
Heat of condensation of VOC: 9.240 Btu/lb-mole
Substitution of these data, the molar flow rates, and the temperatures into
Equations 8.12, 8.15 and 8.16 yields the following enthalpy changes in Btu/hr:
= 130.8Q
D
uncan
The condenser heat load (#;oa(f) is the sum of these three enthalpy changes
(Equation 8.11):
8-47
-------
The refrigeration capacity (#, tons) is computed from Equation 8.22:
This last equation relates the refrigeration capacity (tons) to the inlet gaso-
line vapor flow rate (gal/min). Solving for Qg in terms of R, we obtain:
Qg = 83.9fl
Finally, we substitute this relationship into the equipment cost ($) vs. vapor
flow rate (Qg) correlation, which.was developed from the vendor cost data:
ECp = 58.5Qg + 212,000
= 58.5(83.9fl) + 212, 000
= 4,910^ + 212,000 x
Note that this last expression is identical to Equation 8.34.
8-48
-------
References
[1] Erikson, D.G., Organic Chemical Manufacturing Volume 5: Adsorption,
Condensation, and Absorption Devices, U.S. Environmental Protection
Agency. Research Triangle Park, North Carolina, Publication No. EPA
450/3-80-027, December 1980.
[2] Vatavuk, W.M., and R.B. Neveril, "Estimating Costs of Air Pollution
Control Systems: File Part XVI. Costs of Refrigeration Systems", Chem-
ical Engineering, May 16, 1983, pp. 95-98.
[3] McCabe, W.L., and J.C. Smith, Unit Operations of Chemical Engineer-
ing (Third Edition), McGraw-Hill Book Company, New York, 1976.
[4] Perry, R.H. and C.H. Chilton, Eds. Chemical Engineers' Handbook
(Sixth Edition), McGraw-Hill Book Company, New York, 1989.
[5] Kern, D.Q., Process Heat Transfer, McGraw-Hill Book Company, New
York, 1950.
[6] Smith, J.M., and M.C. VanNess, Introduction to Chemical Engineering
Thermodynamics (Third Edition), McGraw-Hill Book Company, New
York, 1975.
[7] Reid, Robert C., John M. Prausnitz, and Bruce E. Poling, Properties of
Gases & Liquids (Fourth Edition), McGraw-Hill Book Company, New
York, 1987.
r8] Letter and attachment from Robert V. Sisk, Jr. of Piedmont Engineer-
ing, Pineville, North Carolina, to Wiley Barbour of Radian Corporation,
Research Triangle Park, North Carolina, January 28, 1991.
;9j Letter and attachment from Waldrop, R., and V. Sardo of Edwards
Engineering Corp., Pompton Plains, New Jersey, to Wiley Barbour of
8-49
-------
Radian Corporation, Research Triangle Park, North Carolina, October
1, 1990.
[10] Price, Brian C., "Know the Range and Limitations of Screw Compres-
sors", Chemical Engineering Progress, 87(2):50-56.
[11] Letter and attachment from Bob Hansek of ITT Corporation, Atlanta,
Georgia to Wiley Barbour of Radian Corporation, Research Triangle
Park, North Carolina, October 10, 1990.
[12] Letter and attachment from Avery Cooke of Liquid Handling Equip-
ment, Inc., Charlotte, North Carolina to Rich Pelt of Radian Corpora-
tion, Research Triangle Park, North Carolina, September 20, 1990.
[13] Letter from Richard Waldrop of Edwards Engineering Corp., Pompton
Plains, New Jersey to William Vatavuk, P.E., Durham, North Carolina,
August 29, 1988.
[14] Vatavuk, W.M., and R.B. Neveril, "Estimating Costs for Air Pollution
Control Systems, Part II: Factors for Estimating Capital and Operating
Costs", Chemical Engineering, November 3, 1980, pp. 157-162.
[15] Dean, John A., Ed., Lange's Handbook of Chemistry (Twelfth Edition),
McGraw-Hill Book Company, New York, 1978.
8-50
-------
Chapter 9
GAS ABSORBERS
Wiley B arbour
Roy Oommen
Gunseli Sagun Shareef
Radian Corporation
Research Triangle Park, NC 27709
William M. Vatavuk
Standards Development Branch, OAQPS
U. S. Environmental Protection Agency
Research Triangle Park, NC 27711
October 1992
Contents
9.1 Introduction 9.3
9.1.1 System Efficiencies and Performance 9.4
9.2 Process Description 9.5
9.2.1 Absorber System Configuration 9.5
9-1
-------
9.2.2 Types of Absorption Equipment 9.5
9.2.3 Packed Tower Internals 9.7
9.2.4 Packed Tower Operation g.jj
9.3 Design Procedures g_jg
9.3.1 Step 1: Determining Gas and Liquid Stream Conditions 9-14
9.3.2 Step 2: Determining Absorption Factor 9-22
9.3.3 Step 3: Determining Column Diameter 9.23
9.3.4 Step 4: Determining Tower Height and Surface Area . . 9-26
9.3.5 Step '5: Calculating Column Pressure Drop 9-29
9.3.6 Alternative Design Procedure g_29
9.4 Estimating Total Capital Investment 9.31
9.4.1 Equipment Costs for Packed Towers 9.32
9.4.2 Installation Costs q.Qg
9.5 Estimating Annual Cost gL3g
9.5.1 Direct Annual Costs g_3g
9.5.2 Indirect Annual Costs g_40
9.5.3 Total Annual Cost g_4Q
9.6 Example Problem #1 9.41
9.6.1 Required Information for Design 9.41
9.6.2 Step 1: Determine Gas and Liquid Stream Properties . 9-41
9.6.3 Step 2: Calculate Absorption Factor ' 9.45
9.6.4 Step 3: Estimate Column Diameter 9.45
9-2
-------
9.6.5 Step 4: Calculate Column Surface Area 9-47
9.6.6 Step 5: Calculate Pressure Drop 9-48
9.6.7 Equipment Costs 9-48
9.6.8 Total Annual Cost 9-50
9.7 Example Problem #2 9.54
9.8 Acknowledgements 9.54
Appendix 9A - Properties of Pollutants 9-56
Appendix 9B - Packing Characteristics 9.53
Appendix 9C - Minimum Wetting Rate Analysis 9-63
9C.1 Overview of the Approach 9.53
9C.2 Example Problem Calculation 9-64
References 9-66
9.1 Introduction
Gas absorbers are used extensively in industry for separation and purification
of gas streams, as product recovery devices, and as pollution control devices.
This chapter focuses on the application of absorption for pollution control on
gas streams with typical pollutant concentrations ranging from 250 to 10,000
ppmv. Gas absorbers are most widely used to remove water soluble inorganic
contaminants from air streams.[1, 2]
Absorption is a process where one or more soluble components of a gas
mixture are dissolved in a liquid (i.e., a solvent). The absorption process
can be categorized as physical or chemical. Physical absorption occurs when
the absorbed compound dissolves in the solvent; chemical absorption occurs
when the absorbed compound and the solvent react. Liquids commonly used
as solvents include water, mineral oils, nonvolatile hydrocarbon oils, and
aqueous solutions.[1]
9-3
-------
9.1.1 System Efficiencies and Performance
™w «" absorbers ™y for each pollutant-solvent system
and with the type of absorber used. Most absorbers have removal efficiencies
"
a o,
as nigh as 99.9 percent for some pollutant-solvent systems.[l, 3]
The suitability of gas absorption as a pollution control method is gene rally
dependent on the following factors: 1) availability of suitable sofvent; 2^
required removal efficiency; 3) pollutant concentration in the inlet vapor
puuantnt lvent.[l] Absorption is also
enhanced by greater contacting surface, higher liquid-gas ratios and higher
concentrations in the gas stream. [1]
bilitv8114 r11 ^ rem°Ve thC P°Uutant(s) «hould have a high solu-
bility for the gas low vapor pressure, low viscosity, and should be relatively
mexPensive.[4] Water is the most common solvent used to remove inorganic
contammants; lt :s also used to absorb organic compounds having relatively
high water solubihties. For organic compounds that have low water solu
bihties, other solvents such as hydrocarbon oils are used, though only in
mdustnes where large volumes of these oils are available (i.e.,
refineries and petrochemical plants).[5]
of the0!4"11!-'61110?! may ^ bC enhanced ^ manipulating the chemistry
of the absorbing ,olut,on so that it reacts with the pollutant(s), e.^ caustic
solution for aad-ga. absorption vs. pure water as a solvent. Chemical ab-
orption may be hmited by the rate of reaction, although the rate limiting
step ,. typlcally the physical absorption rate, not the chemical reaction rate
9-4
-------
9.2 Process Description
Absorption is a mass transfer operation in which one or more soluble com-
ponents of a gas mixture are dissolved in a liquid that has low volatility
under the process conditions. The pollutant diffuses from the gas into the
liquid when the Uquid contains less than the equilibrium concentration of the
gaseous component. The difference between the actual concentration and the
equilibrium concentration provides the driving force for absorption.
A properly designed gas absorber will provide thorough contact between
the gas and the solvent in order to facilitate diffusion of the pollutant(s).
It will perform much better than a poorly designed absorber.[6] The rate
of mass transfer between the two phases is largely dependent on the surface
area exposed and the time of contact. Other factors governing the absorption
rate, such as the solubility of the gas in the particular solvent and the degree
of the chemical reaction, are characteristic of the constituents involved and
are relatively independent of the equipment used.
9.2.1 Absorber System Configuration
Gas and liquid flow through an absorber may be countercurrent, cross-
current, or cocurrent. The most commonly installed designs are countercur-
rent, in which the waste gas stream enters at the bottom of the absorber col-
umn and exits at the top. Conversely, the solvent stream enters at the top and
exits at the bottom. Countercurrent designs provide the highest theoretical
removal efficiency because gas with the lowest pollutant concentration con-
tacts liquid with the lowest pollutant concentration. This serves to maximize
the average driving force for absorption throughout the column. [2] Moreover
countercurrent designs usually require lower liquid to gas ratios than cocur-
rent and are more suitable when the pollutant loading is higher.[3, 5]
In a crosscurrent tower, the waste gas flows horizontally across the column
while the solvent flows vertically down the column. As a rule, crosscurrent
designs have lower pressure drops and require lower liquid-to-gas ratios than
both cocurrent and countercurrent designs. They are applicable ^hen
-------
the top of the tower and exit at the bottom. Cocurrent designs have lower
pressure drops, are not subject to flooding limitations and are more efficient
for fine (t.e., submicron) mist removal. Cocurrent designs are only efficient
where large absorption driving forces are available. Removal efficiency is
limited since the gas-liquid system approaches equilibrium at the bottom of
the tower. [2]
9.2.2 Types of Absorption Equipment
Devices that are based on absorption principles include packed towers, plate
(or tray) columns, venturi scrubbers, and spray chambers. This chapter fo-
cuses on packed towers, which are the most commonly used gas absorbers for
pollution control. Packed towers are columns filled with packing materials
that provide a large surface area to facilitate contact between the liquid and
gas. Packed tower absorbers can achieve higher removal efficiencies, handle
higher liquid rates, and have relatively lower water consumption requirements
than other types of gas absorbers.[2] However, packed towers may also have
high system pressure drops, high clogging and fouling potential, and exten-
sive maintenance costs due to the presence of packing materials. Installation,
operation, and wastewater disposal costs may also be higher for packed bed
absorbers than for other absorbers.[2] In addition to pump and fan power re-
quirements and solvent costs, packed towers have operating costs associated
with replacing damaged packing.[2]
»
Plate, or tray, towers are vertical cylinders in which the liquid and gas
are contacted in step-wise fashion on trays (plates). Liquid enters at the
top of the column and flows across each plate and through a downspout
(downcomer) to the plates below. Gas moves upwards through openings in
the plates, bubbles into the liquid, and passes to the plate above. Plate
towers are easier to clean and tend to handle large temperature fluctuations
better than packed towers do.[4] However, at high gas flow rates, plate towers
exhibit larger pressure drops and have larger liquid holdups. Plate towers
are generally made of materials such as stainless steel, that can withstand
the force of the liquid on the plates and also provide corrosion protection.
Packed columns are preferred to plate towers when acids and other corrosive
materials are involved because tower construction can then be of fiberglass,
polyvinylchloride, or other less costly, corrosive-resistant materials. Packed
towers are also preferred for columns smaller than two feet in diameter arid
9-6
-------
when pressure drop is an important consideration. [3, 7]
Venturi scrubbers are generally applied for controlling particulate mat-
ter and sulfur dioxide. They are designed for applications requiring high
removal efficiencies of submicron particles, between 0.5 and 5.0 micrometers
in diameter. [4] A venturi scrubber employs a gradually converging and then
diverging section, called the throat, to clean incoming gaseous streams. Liq-
uid is either introduced to the venturi upstream of the throat or injected
directly into the throat where it is atomized by the gaseous stream. Once
the liquid is atomized, it collects particles from the gas and discharges from
the venturi.[1] The high pressure drop through these systems results in high
energy use, and the relatively short gas-liquid contact time restricts their
application to highly soluble gases. Therefore, they are infrequently used for
the control of volatile organic compound emissions in dilute concentration. [2]
Spray towers operate by delivering liquid droplets through a spray dis-
tribution system. The droplets fall through a countercurrent gas stream un-
der the influence of gravity and contact the pollutant(s) in the gas.[7] Spray
towers are simple to operate and maintain, and have relatively low energy
requirements. However, they have the least effective mass transfer capability
of the absorbers discussed and are usually restricted to particulate removal
and control of highly soluble gases such as sulfur dioxide and ammonia. They
also require higher water recirculation rates and are inefficient at removing
very small particles.[2, 5]
9.2.3 Packed Tower Internals
A basic packed tower unit is comprised of a column shell, mist eliminator,
liquid distributors, packing materials, packing support, and may include a
packing restrainer. Corrosion resistant alloys or plastic materials such as
polypropylene are required for column internals when highly corrosive sol-
vents or gases are used. A schematic drawing of a countercurrent packed
tower is shown in Figure 9.1. In this figure, the packing is separated into
two sections. This configuration is more expensive than designs where the
packing is not so divided. [5]
The tower shell may be made of steel or plastic, or a combination of these
materials depending on the corrosiveness of the gas and liquid streams, and
the process operating conditions. One alloy that is chemical and temperature
9-7
-------
f QuOut
Uquid In —
Mitt Eliminator
U'quid Distributor
Spray Nozzt*
Packing
Picking
Liquid R*-distnbutor
Picking Support
Gttln
Uquid Out
Figure 9.1: Packed Tower for Gas Absorption
9-8
-------
resistant or multiple layers of different, less expensive materials may be used.
The shell is sometimes lined with a protective membrane, often made from a
corrosion resistant polymer. For absorption involving acid gases, an interior
layer of acid resistant brick provides additional chemical and temperature
resistance. [8]
At high gas velocities, the gas exiting the top of the column may carry
off droplets of liquid as a mist. To prevent this, a mist eliminator in the
form of corrugated sheets or a layer of mesh can be installed at the top of
the column to collect the liquid droplets, which coalesce and fall back into
the column.
A liquid distributor is designed to wet the packing bed evenly and
initiate uniform contact between the liquid and vapor. The liquid distributor
must spread the liquid uniformly, resist plugging and fouling, provide free
space for gas flow, and allow operating flexibility.[9] Large towers frequently
have a liquid redistributor to collect liquid off the column wall and direct it
toward the center of the column for redistribution and enhanced contact in
the lower section of packing. [4] Liquid redistributors are generally required
for every 8 to 20 feet of random packing depth.[5, 10]
Distributors fall into two categories: gravitational types, such as orifice
and weir types, and pressure-drop types, such as spray nozzles and perfo-
rated pipes. Spray nozzles are the most common distributors, but they may
produce a fine mist that is'easily entrained in the gas flow. They also may
plug, and usually require high feed rates to compensate for poor distribution.
Orifice-type distributors typically consist of flat trays with a number of risers
for vapor flow and perforations in the tray floor for liquid flow. The trays
themselves may present a resistance to gas flow.[9] However, better contact
is generally achieved when orifice distributors are used. [3]
Packing materials provide a large wetted surface for the gas stream
maximizing the area available for mass transfer. Packing materials are avail-
able in a variety of forms, each having specific characteristics with respect to
surface area, pressure drop, weight, corrosion resistance, and cost. Packing
life varies depending on the application. In ideal circumstances, packing will
last as long as the tower itself. In adverse environments packing life may be
as. short as 1 to 5 years due to corrosion, fouling, and breakage".[11]
Packing materials are categorized as random or structured. Random
packings are usually dumped into an absorption column and allowed to.settle.
9-9
-------
Pall Ring
Tellerecte
Incalox Saddle
Berl Saddle
Raschig Ring
Figure 9.2: Random Packing Materials
Modern random packings consist of engineered shapes intended to maximize
surface-to-volume ratio and minimize pressure drop.[2] Examples of different
random packings are presented in Figure 9.2. The first random packings
specifically designed for absorption towers were made of ceramic. The use of
ceramic has declined because of their brittleness, and the current markets are
dominated by metal and plastic. Metal packings cannot be used for highly
corrosive pollutants, such as acid gas, and plastic packings are not suitable
for high temperature applications. Both plastic and metal packings are gen-
erally limited to an unsupported depth of 20 to 25. At higher depths the
weight may deform the packing. [10]
Structured packing may be random packings connected in an orderly
arrangement, interlocking grids, or knitted or woven wire screen shaped
into cylinders or gauze like arrangements. They usually have smaller pres-
sure drops and are able to handle greater solvent flow rates than random
9-10
-------
packings.[4] However, structured packings are more costly to install and may
not be practical for smaller columns. Most structured packings are made
from metal or plastic.
In order to ensure that the waste gas is well distributed, an open space
between the bottom of the tower and the packing is necessary. Support
plates hold the packing above the open space. The support plates must
have enough strength to carry the weight of the packing, and enough free
area to allow solvent and gas to flow with minimum restrictions.[4]
High gas velocities can fluidize packing on top of a bed. The packing
could then be carried into the distributor, become unlevel, or be damaged.[9]
A packing restrainer may be installed at the top of the packed bed to
contain the packing. The packing restrainer may be secured to the wall so
that column upsets will not dislocate it, or a "floating" unattached weighted
plate may be placed on top of the packing so that it can settle with the bed.
The latter is often used for fragile ceramic packing.
9.2.4 Packed Tower Operation
As discussed in Section 9.2.1, the most common packed tower designs are
countercurrent. As the waste gas flows up the packed column it will experi-
ence a drop in its pressure as it meets resistance from the packing materials
and the solvent flowing down. Pressure drop in a column is a function of
the gas and liquid flow rates and properties of the packing elements, such as
surface area and free volume in the tower. A high pressure drop results in
high fan power to drive the gas through the packed tower, and consequently
high costs. The pressure drop in a packed tower generally ranges from 0 5 to
1.0 in. H2O/ft of packing.[7]
For each column, there are upper and lower limits to solvent and vapor
flow rates that ensure satisfactory performance. The gas flow rate may be-
come so high that the drag on the solvent is sufficient to keep the solvent from
flowing freely down the column. Solvent begins to accumulate and blocks the
entire cross section for flow, which increases the pressure drop and prevents
the packing from mixing the gas and solvent effectively. When all the free
volume in the packing is filled with liquid and the liquid is carried back up
the column, the absorber is considered to be flooded.[4| Most packed towers
operate at 60 to 70 percent of the gas flooding velocity," as it is not practical to
9-11
-------
operate a tower in a flooded condition.^] A minimum liquid flow rate is also
required to wet the packing material sufficiently for effective mass transfer
to occur between the gas and liquid.[7]
The waste gas inlet temperature is another important scrubbing param-
eter. In general, the higher the gas temperature, the lower the absorption
rate, and vice-versa. Excessively high gas temperatures also can lead to
significant solvent loss through evaporation. Consequently, precoolers (e.g.
spray chambers) may be needed to reduce the air temperature to acceptable
levels. [6]
For operations that are based on chemical reaction with absorption, an
additional concern is the rate of reaction between the solvent and pollu-
tant^). Most gas absorption chemical reactions are relatively fast and the
rate limiting step is the physical absorption of the pollutant(s) into the sol-
vent. However, for solvent-pollutant systems where the chemical reaction is
the limiting step, the rates of reaction would need to be analyzed kinetically.
Heat may be generated as a result of exothermal chemical reactions. Heat
may also be generated when large amounts of solute are absorbed into the
liquid phase, due to the heat of solution. The resulting change in temper-
ature along the height of the absorber column may damage equipment and
reduce absorption efficiency. This problem can be avoided by adding cooling
coils to the column.[7] However, in those systems where water is the solvent,
adiabatic saturation of the gas occurs during absorption due to solvent evap-
oration. This causes a substantial cooling of the absorber that offsets the
heat generated by chemical reactions. Thus, cooling coils are rarely required
with those systems.[5] In any event, packed towers may be designed assuming
that isothermal conditions exist throughout the column.[7]
The effluent from the column may be recycled into the system and used
again. This is usually the case if the solvent is costly, i.e., hydrocarbon oils,
caustic solution. Initially, the recycle stream may go to a waste treatment
system to remove the pollutant(s) or the reaction product. Make-up solvent
may then be added before the Uquid stream reenters the column. Recircula-
tion of the solvent requires a pump, solvent recovery system, solvent holding
and mixing tanks, and any associated piping and instrumentation.
9-12
-------
9.3 Design Procedures
The design of packed tower absorbers for controlling gas streams containing
a mixture of pollutant(s) and air depends on knowledge of the Mowing
parameters:
1. Waste gas flow rate;
2. Waste gas composition and concentration of the pollutant(s) in the gas
stream;
3. Required removal efficiency;
4. Equilibrium relationship between the pollutant(s) and solvent; and
5. Properties of the pollutant(s), waste gas, and solvent:
• Diffusivity,
• Viscosity,
• Density, and
• Molecular weight.
The primary objectives of the design procedures are to determine column
surface area and pressure drop through the column. In order to determine
these parameters, the following steps must be performed:
Step 1: Determine the gas and liquid stream conditions entering and exiting
the column.
Step 2: Determine the absorption factor
Step 3: Determine the diameter of the column (Z>).
Step 4: Determine the tower height (fftower) and surface area (5).
Step 5: Determine the packed column pressure drop (AP).
9-13
-------
To simplify the sizing procedures, a number of assumptions have been
made. For example, the waste gas is assumed to comprise a two-component
waste gas mixture (pollutant/air), where the pollutant consists of a single
compound present in dilute quantities. The waste gas is assumed to behave
as an ideal gas and the solvent is assumed to behave as an ideal solution.
Heat effects associated with absorption are considered to be minimal for the
pollutant concentrations encountered. The procedures also assume that, in
chemical absorption, the process is not reaction rate limited, t.e., the reaction
of the pollutant with the solvent is considered fast compared to the rate of
absorption of the pollutant into the solvent.
The design procedures presented here are complicated, and careful atten-
tion to units is required. Table 9.1 is a list of all design variables referred
to in this chapter, along with the appropriate units. A key is provided to
differentiate primary data from calculated data.
9.3.1 Step 1: Determining Gas and Liquid Stream
Conditions
Gas absorbers are designed based on the ratio of liquid to gas entering the
column (!,/
-------
Table 9.1: List of Design Variables
Variable
Symbol
Units
> Surface to volume ratio of
packing
Cross-sectional area of ab-
sorber
Abscissa value from plot of
generalized pressure drop cor-
relation
Absorption factor
Diameter of absorber
> Diffusivity of pollutant in gas
> Diffusivity of pollutant in liq-
uid
> Flooding factor
t> Packing factor
> Waste gas flow rate entering
absorber
Waste gas flow rate exiting ab-
sorber
Waste gas molar flow rate en-
tering absorber
Molar flow rate of pollutant
free gas
Waste gas superficial flow rate
entering absorber
Height of gas transfer unit
Height of liquid transfer unit
Height of overall transfer unit
Height of packing
Height of absorber
Pressure drop constants
Liquid rate entering absorber
Liquid rate exiting absorber
Liquid molar flow rate enter-
ing absorber
Molar flow rate of pollutant
- free solvent
a
A
ABSCISSA
AF
D
DG
DL
f
G]
G
mol
G,
G
,fr,i
HI
H
TT
pack
tower
L,
'mol,i
ft2/ft:3
ft2
feet
ft2/hr
ft2/hr
acfm
acfm
Ib-moles/hr
Ib-moles/hr
lb/sec-ft2
feet
feet
feet
feet
feet
gpm
gpm
Ib-moles/hr
Ib-moles/hr
9-15
-------
Table 9.1: List of Design Variables (Continued)
Variable
Symbol
Units
Liquid superficial flow rate en-
tering absorber
Slope of equilibrium line
t> Molecular weight of the gas
stream
t> Molecular weight of the liquid
stream
> Minimum wetting rate
Number of overall transfer
units
Ordinate value from plot of
generalized pressure drop cor-
relation
Surface area of absorber
t> Temperature of solvent
Mole fraction of pollutant en-
tering absorber in liquid
Mole fraction of pollutant ex-
iting absorber in liquid
Pollutant concentration enter-
ing absorber in liquid
Maximum pollutant concen-
tration in liquid phase in equi-
librium with pollutant enter-
ing column in gas phase
Pollutant concentration exit-
•ing absorber in liquid
Mole fraction of pollutant en-
tering absorber in waste gas
Mole fraction of pollutant in
gas phase in equilibrium with
mole fraction of pollutant en-
tering in the liquid phase
Mole fraction of pollutant ex-
iting scrubber in waste gas
m
MWG
MWL
MWR
*f.
ORDINATE
5
T
lb/hr-ft2
Ib/lb-mole
Ib/lb-mole
ft2/hr
ft2
K
Ib-mole of pollutant
Ib-mole totaTliquid
Ib-mole of pollutant
Ib-mole total liquid
Ib-moles pollutant
Ib-moles pollutant free solvent
Ib-moles pollutant
Ib-moles pollutant tree solvent
Ib-moles pollutant
Ib-moles pollutant free solvent
Ib-moles pollutant
Ib-moles total gasi
Ib-moles pollutant
Ib-moles total gas
Ib-moles pollutant
Ib-moles total gas
9-16
-------
Table 9.1: List of Design Variables (Continued)
Variable
Symbol
Units
Mole fraction of pollutant in
gas phase in equilibrium with
mole fraction of pollutant ex-
iting in the liquid phase
> Pollutant concentration enter-
ing scrubber in waste gas
Pollutant concentration enter-
ing scrubber in equilibrium
with concentration in liquid
phase
Pollutant concentration exit-
ing scrubber in waste gas
Pollutant concentration exit-
ing scrubber in equilibrium
with concentration in liquid
phase
t> Pollutant removal efficiency
> Density of waste gas stream
t> Density of liquid stream
C> Viscosity of waste gas
> Viscosity of solvent
Ratio of solvent density to wa-
ter density
Pressure drop
> Packing factors
t> Denotes required input data.
Y*
PC
PL
V-G
fJ-r,
*
AP
Ib-moles pollutant
Ib-moles total gas
Ib-moles pollutant
ib-moles pollutant free gas
Ib-moles pollutant___
Ib-moles pollutant free gas
Ib-moles pollutant
Ib-moles pollutant free gas
Ib-moles pollutant
Ib-moles pollutant free gas
lb/ft3
lb/ft3
Ib/ft-hr
Ib/ft-hr
inches H20/feet of packing
9-17
-------
This design approach assumes that the inlet gas stream variables are
known, and that a specific pollutant removal efficiency has been chosen as
the design basis; t.e., the variables (?,-, Yh and 77 are known. For dilute
concentrations typically encountered in pollution control applications and
negligible changes in moisture content, G, is assumed equal to G0. If a
once-through process is used, or if the spent solvent is regenerated by an air
stripping process before it is recycled, the value of X{ will approach zero.
The Mowing procedures must be Mowed to calculate the remaining stream
variables Y0, I, (and L0), and X0. A schematic diagram of a packed tower
with inlet and outlet flow and concentration variables labeled is presented in
Figure 9.3.
The variable Y0 may be calculated from 77 using the Mowing equation:
Y-= Y- (' - is) (••!)
The liquid flow rate entering the absorber, Z, (gpm), is then calculated
using a graphical method. Figure 9.4 presents an example of an equilibrium
curve and operating line. The equilibrium curve indicates the relationship
between the concentration of pollutant in the waste gas and the concentra-
tion of pollutant in the solvent at a specified temperature. The operating line
indicates the relation between the concentration of the pollutant in the gas
and solvent at any location in the gas absorber column. The vertical distance
between the operating line and equilibrium curve indicates the driving force
for diffusion of the pollutant between the gas and liquid phases. The mini-
mum amount of liquid which can be used to absorb the pollutant in the gas
stream corresponds to an operating line drawn from the outlet concentration
m the gas stream (F0) and the inlet concentration in the solvent stream (X,)
to the point on the equilibrium curve corresponding to the entering pollu-
tant concentration in the gas stream (Y,). At the intersection point on the
equilibrium curve, the diffusional driving forces are zero, the required time
of contact for the concentration change Is infinite, and an infinitely tall tower
results.
The slope of the operating line intersecting the equilibrium curve is equal
to the minimum L/G ratio on a moles of pollutant-free solvent (L,) per moles
of pollutant-free gas basis (G,). In other words, the values L, and Gfdo not
include the moles of pollutant in the liquid and gas streams. The values of £,
and G, are constant through the column if a negligible amount of moisture
is transferred from the liquid to the gas phase. The slope may be calculated
9-18
-------
GO
gmol, o
yS
v°
y o
YT
y I
mol, o
Figure 9.3: Schematic Diagram of Countercurrent Packed Tower Operati
on
9-19
-------
8
d)
"o
Q.
"5
Yfl .._
Moles of Pollutant/Mole of Solvent
Figure 9.4: Minimum and Actual Liquid-to-Gas Ra.il
OS
9-20
-------
from the following equation:
(L>\ Y'~YO
\G~) • = "F — F (9-2)
Xl-rs/ mtn -*-0 —A, '
where X* would be the maximum concentration of the pollutant in the liquid
phase if it were allowed to come to equilibrium with the pollutant entering the
column in the gas phase, 1-. The value of X'0 is taken from the equilibrium
curve. Because the minimum L./G. ratio is an unrealistic value, it must
be multiplied by an adjustment factor, commonly between 1.2 and 1.5, to
calculate the actual L/G ratio:[7]
7^~) ~ \7^~} x (adjustment factor) (9.3)
<->•*/ ^/' '
act
The variable G, may be calculated using the equation:
_ 60 pGGi
G° ~ MW0(1 + Yi) (9-4)
where 60 is the conversion factor from minutes to hours, MWG is the molec-
ular weight of the gas stream (Ib/lb-mole), and pG is the density of the
gas stream (lb/ftj). For pollutant concentrations typically encountered, the
molecular weight and density of the waste gas stream are assumed to be
equal to that of ambient air.
The variable Ls may then be calculated by:
The total molar flow rates of the gas and liquid entering the absorber (Gmol •
and Lmol,i) are calculated using the following equations:
=»l + t (9.7)
The volume flow rate of the solvent, I,, may then be calculated by using the
following relationship:
where 60 is the conversion factor from minutes to hours, MW/ is the molec-
ular weight of the liquid stream (Ib/lb-moie), pL is the density of the liquid
9-21
-------
stream (lb/ft3), and 7.48 is the factor used to convert cubic feet to gallons.
If the volume change in the liquid stream entering and exiting the absorber
is assumed to be negligible, then X, = L0.
Gas absorber vendors have provided a range for the £,/(?, ratio for acid
gas control from 2 to 20 gpm of solvent per 1000 cfm of waste gas. [12] Even
for pollutants that are highly soluble in a solvent (i.e., HC1 in water), the
adjusted Li/Gi ratio calculated using Equations 9.2 to 9.8 would be much
lower than this range, because these equations do not consider the flow rate
of the solvent required to wet the packing.
Finally, the actual operating line may be represented by a material bal-
ance equation over the gas absorber: [4]
XiL, + YiG, = X0L. + Y0G, (9.9)
Equation 9.9 may then be solved for X0:
x° =
9.3.2 Step 2: Determining Absorption Factor
The absorption factor (AF) value is frequently used to describe the rela-
tionship between the equilibrium line and the liquid-to-gas ratio. For many
pollutant-solvent systems, the most economical value for AF ranges around
1.5 to 2.0. [7] The Mowing equation may be used to calculate AF: [4, 7]
A j-, "mol.i
AF = m r 9.11)
m Gmol,i V '
where m is the slope of the equilibrium line on a mole fraction basis. The
value of m may be obtained from available literature on vapor/liquid equilib-
rium data for specific systems. Since the equilibrium curve is typically linear
in the concentration ranges usually encountered in air pollution control, the
slope, m would be constant (or nearly so) for all applicable inlet and out-
let liquid and gas streams. The slope may be calculated from mole fraction
values using the following equation: [4]
™ = rr7 (9.12)
U>0 .i/j
9-22
-------
where yj and y'0 are the mole fractions of the pollutant in the vapor phase in
equilibrium with the mole fractions of the pollutant entering and exiting the
absorber in the liquid, z, and z0, respectively. The slope of the equilibrium
line in Figure 9.4 is expressed in terms of concentration values ^Y,, X0, Yf,
and Y*. These values may be converted to x,, x0, y?, and y; using the
equations:
(9.13)
(9.14)
(9.15)
(9.16)
where the units for each of these variables are listed in Table 9.1.
The absorption factor will be used to calculate the theoretical number of
transfer units and the theoretical height of a transfer unit. First, however,
the column diameter needs to be determined.
9.3.3 Step 3: Determining Column Diameter
Once stream conditions have been determined, the diameter of the column
may be estimated. The design presented in this section is based on select-
ing a fraction of the gas flow rate at flooding conditions. Alternatively, the
column may be designed for a specific pressure drop (see Section 9.3.6.). Eck-
ert's modification to the generalized correlation for randomly packed towers
based on flooding considerations is used to obtain the superficial gas flow
rate entering the absorber, Gafr>i (lb/sec-ft2), or the gas flow rate per cross-
sectional area based on the Lmo[>i/Gmolji ratio calculated in Step 2.[10] The
cross-sectional area (.4) of the column and the column diameter (D) can then
be determined from (?„/,.,. - •
sjr, i
Figure 9.5 presents the relationship between Gsfr i and the Lmol -,/G
ratio at the tower flood point. The abscissa value (X axis) in the graph"is
9-23
-------
Figure 9.5: Eckert's Modiilcation to the Generalized Correlation at Flood-
ing .Rate[10]
9-24
-------
expressed as:[10]
ABSCISSA =
Gmol,
The ordinate value (Y axis) in the graph is expressed as:[10]
(9.18)
I u/* I r~u k°'2
ORDINATE = "" " '' w* '^
PLRG9c
where Fp is a packing factor, gc is the gravitational constant (32.2), /J.L is the
viscosity of the solvent (Ib/ft-hr), 2.42 is the factor used to convert Ib/ft-hr to
centipoise, and * is the ratio of the density of the scrubbing Uquid to water.
The value of Fp may be obtained from packing vendors (see Appendix 9B
Table 9.8).
After calculating the ABSCISSA value, a corresponding ORDINATE
value may determined off the flooding curve. The ORDINATE may also
be calculated using the following equation:[10]
ORDINATE = io[-1-868-I-°w
-------
If a substantial change occurs between inlet and outlet volumes (i.e., moisture
is transferred from the liquid phase to the gas phase), the diameter of the
column will need to be calculated at the top and bottom of the column. The
larger of the two values is then chosen as a conservative number. As a, rule
of thumb, the diameter of the column should be at least 15 times the size of
the packing used in the column. If this is not the case, the column diameter
should be recalculated using a smaller diameter packing. [10]
The superficial liquid flow rate entering the absorber, Lafri (lb/hr-ft2),
ba.sed on the cross- sectional area determined in Equation 9.21 'is calculated
from the equation:
r Lmol,i
For the absorber to operate properly, the liquid flow rate entering the
column must be high enough to effectively wet the packing so mass transfer
between the gas and liquid can occur. The minimum value of Lsfr t that is
required to wet the packing effectively can be calculated using the equation- \7
13] "l '
(L^min = MWRP^ (9.24)
where M WR is defined as the minimum wetting rate (ft2/hr), and a is the
surface area to volume ratio of the packing (ft2/ft3). An MWR value of
0.85 ft2/hr js recommended for ring packings larger than 3 inches and for
structured grid packings. For other packings, an MWR of 1.3 ft2/hr is
recommended. [7, 13] Appendix 9B, Table 9.8 contains values of a for common
packing materials.
If Ltfr,i (the value calculated in Equation 9.23) is smaller than (Lafr i}min
(the value calculated in Equation 9.24), there is insufficient liquid flow to wet
the packing using the current design parameters. The value of G ^ ,, and A
then will need to be recalculated. See Appendix 9C for details. * '*
9.3.4 Step 4: Determining Tower Height and Surface
Area
Tower height is primarily a function of packing depth. The required depth of
packing (Hpack) is determined from the theoretical number of overall transfer
9-26
-------
units (Ntu) needed to achieve a specific removal efficiency, and the height of
the overall transfer unit (Htu):[4]
Hpack = Ntu Htu (9.25)
The number of overall transfer units may be estimated graphically by step-
ping off stages on the equilibrium-operating line graph from inlet conditions
to outlet conditions, or by the following equation:[4]
(9.26)
AF
where In is the natural logarithm of the quantity indicated. The equation
is based on several assumptions: 1) Henry's law applies for a dilute gas
mixture; 2) the equilibrium curve is linear from x,- to xa; and 3) the pollutant
concentration in the solvent is dilute enough such that the operating line can
be considered a straight line. [4]
If z, -> 0 (t'.e., a negligible amount of pollutant enters the absorber in the
liquid stream) and l/AF -> 0 (t.e., the slope of the equilibrium line is very
small and/or the Lmol/Gmol ratio is very large), Equation 9.26 simplifies to:
(9.27)
There are several methods that may be used to calculate the height of
the overall transfer unit, all based on empirically determined packing con-
stants. One commonly used method involves determining the overall gas
and liquid mass transfer coefficients (Kc, KL). A major difficulty in using
this approach is that values for Kc and KL are frequently unavailable for
the specific pollutant-solvent systems of interest. The reader is referred to
the book Random Packing and Packed Tower Design Applications in the
reference section for further details regarding this method.[14]
For this chapter, the method used to calculate the height of the overall
transfer unit is based on estimating the height of the gas and liquid film
transfer units, HL and HG, respectively:[4]
Htu = HG + -r^HL (9.28)
AF
9-27
-------
The Mowing correlations may be used to estimate values for HL and
HG:[13\
(9.29)
(9.30)
The quantity fi/pD is the Schmidt number and the variables a, 0, 7 0
and b are packing constants specific to each packing type. Typical values'for
these constants are listed in Appendix 9B, Tables 9.9 and 9.10. The advan-
tage to using this estimation method is that the packing constants may be
applied to any pollutant-solvent system. One packing vendor offers the fol-
lowing modifications to Equations 9.29 and 9.30 for their specific packing:[15]
Hc,=
01
(9.31)
-4.255
(9.32)
where T is the temperature of the solvent in Kelvin.
After solving for Hpack using Equation 9.25, the total height of the column
may be calculated from the following correlation:[16]
Htower = 1-40 Hpack + 1.02 D + 2.81
(9.33)
Equation 9.33 was developed from information reported by gas absorber ven-
dors, and is applicable for column diameters from 2 to 12 feet and packing
depths from 4 to 12 feet. The surface area (S) of the gas absorber can be
calculated using the equation: [16]
(9.34)
Equation 9.34 assumes the ends of the absorber are flat and circular.
9-28
-------
9.3.5 Step 5: Calculating Column Pressure Drop
Pressure drop in a gas absorber is a function of Gtfrii and properties of the
packing used. The pressure drop in packed columns' generally ranges from
0.5 to 1 inch of H20 per foot of packing. The absorber may be designed
for a specific pressure drop or pressure drop may be estimated using Leva's
correlation • F7 1 ftl
correlation: [7, 10]
PG
The packing constants c and j are found in Appendix 9B, Table 9.11, and
3600 is the conversion factor from seconds to hours. The equation was orig-
inally developed for air-water systems. For other liquids, Lsfr { is multiplied
by the ratio of the density of water to the density of the liquid.
9.3.6 Alternative Design Procedure
The diameter of a column can be designed for a specific pressure drop, rather
than being determined based on a fraction of the flooding rate. Figure 9.6
presents a set of generalized correlations at various pressure drop design
values. The abscissa value of the graph is similar to Equation 9.17:[10]
ABSCISSA =
The ordinate value is expressed as:[10]
ORDINATE =
For a calculated ABSCISSA value, a corresponding ORDINATE value at
each pressure drop can be read off Figure 9.6 or can be calculated from the
following equation:[10]
ORDINATE = exp [*,, + *,(ln ABSCISSA) + k,(ln ABSCISS A)2 +
fc,(ln ABSCISSA) ' + jfe,(hi ABSCISSA) '] (9.38)
The constants fc,,, *„ fc2, &,, and k, are shown below for each pressure drop
value. r
9-29
-------
$
.J
Q.
at
13
AP-15
Figure 9.6: Generalized Pressure Drop CorrelationsflO]
9-30
-------
Constants for Each Pressure Drop Correlation
(inches water/
ft packing)
0.05
0.10
0.25
0.50
1.00
1.50
, , , , ,
-6.3025
-5.5009
-5.0032
-4.3992
-4.0950
-4.0256
-0.6080
-0.7851
-0.9530
-0.9940
-1.0012
-0.9895
-0.1193
-0.1350
-0.1393
-0.1698
-0.1587
-0.0830
-0.0068
0.0013
0.0126
0.0087
0.0080
0.0324
0.0003
0.0017
0.0033
0.0034
0.0032
0.0053
Equation 9.37 can be solved for
G,fr,i =
(PL - pc)p(7(7c(ORDINATE)
O.l
(9.39)
The remaining calculations to estimate the column diameter and Lsfr t are the
same as presented in Section 9.3.3, except the flooding factor (/) is not used
in the equations. The flooding factor is not required because an allowable
pressure drop that will not cause flooding is chosen to calculate the diameter
rather than designing the diameter at flooding conditions and then taking a
fraction of that value.
9.4 Estimating Total Capital Investment
This section presents the procedures and data necessary for estimating cap-
ital costs for vertical packed bed gas absorbers using countercurrent flow
to remove gaseous pollutants from waste gas streams. Equipment costs for
packed bed absorbers are presented in Section 9.4.1, with installation costs
presented in Section 9.4.2.
Total capital investment, TCI, includes equipment cost, EC, for the entire
^ absorber unit, taxes, freight charges, instrumentation, and direct and
indirect installation costs. All costs are presented in third quarter 1991
dollars. The costs presented are study estimates with an expected accuracy
of ± 30 percent. It must be kept in mind that even for a given application,
9-31
-------
design and manufacturing procedures vary from vendor to vendor, so costs
vary. All costs are for new plant installations; no retrofit cost considerations
are included.
9.4.1 Equipment Costs for Packed Towers
Gas absorber vendors were asked to supply cost estimates for a range of
tower dimensions (i.e., height, diameter) to account for the varying needs
of different applications. The equipment for which they were asked to pro-
vide costs consisted of a packed tower absorber made of fiberglass reinforced
plastic (FRP), and to include the Mowing equipment components:
• absorption column shell;
• gas inlet and outlet ports;
• liquid inlet port and outlet port/drain;
• liquid distributor and redistributor;
• two packing support plates;
• mist eliminator;
• internal piping;
• sump space; and
• platforms and ladders.
The cost data the vendors supplied were first adjusted to put them on a
common basis, and then were regressed against the absorber surface area (.5).
The equation shown below is a multivariant regression of cost data provided
by six vendors.[16, 12]
Total Tower Cost($) = 115 S . (9.40)
where S is the surface area of the absorber, in ft2.
Figure 9.7 depicts a plot of Equation 9.40. This equation is applicable
for towers with surface areas from 69 to 1507 ft2 constructed of FRP. Costs
9-32
-------
200,000
180.000
160,000
140,000
~ 120.000
a
3
S! 100,000
$
O 80,000
o-
Ul
60,000
40,000
20.000
200 400 600 800 1,000
Surface Area of Tower (tt2)
1.200
1,400
1.600
Figure 9.7: Packed Tower Equipment Cost[16j
9-33
-------
for towers made of materials other than FRP may be estimated using the
following equation:
TTCA/ = CF x TTC (9.41)
where TTCM is the total cost of the tower using other materials, and TTC
is the total tower cost as estimated using Equation 9.40. The variable CF
is a cost factor to convert the cost of an FRP gas absorber to an absorber
fabricated from another material. Ranges of cost factors provided by vendors
are listed for the following materials of construction:[12]
304 Stainless steel = 1.10-1.75
Polypropylene = 0.80 - 1.10
Polyvinyl chloride = 0.50 - 0.90
Auxiliary costs encompass the cost of all necessary equipment not in-
cluded in the absorption column unit. Auxiliary equipment includes packing
material, instruments and controls, pumps, and fans. Cost ranges for various
types of random packings are presented in Table 9.2. The cost of structured
packings varies over a much wider range. Structured packings made of stain-
less steel range from $45/ft3 to $405/ft3, and those made of polypropylene
range from $65/ft:i to $350/ft;'.[17]
Similarly, the cost of instruments and controls varies widely depending
on the complexity required. Gas absorber vendors have provided estimates
ranging from $1,000 to $10,000 per column. A factor of 10 percent of the
tower cost will be used to estimate this cost in this chapter. Design and cost
correlations for fans and pumps will be presented in a chapter on auxiliary
equipment elsewhere in this manual. However, cost data for auxiliaries are
available from the literature (see reference [18], for example).
The total equipment cost (EC) is the sum of the component equipment
costs, which includes tower cost and the auxiliary equipment cost.
EC = TTC + Packing Cost + Auxiliary Equipment (9.42)
The purchased equipment cost (PEC) includes the cost of the absorber
with packing and its auxiliaries (EC), instrumentation (0.10 EC), sales tax
(0.03 EC), and freight (0.05 EC). The PEC is calculated from the following
factors, presented in Chapter 2 of this manual and confirmed from the gas
absorber vendor survey conducted during this study:[l2, 19]
PEC - (1 + 0.10 + 0.03 + 0.05) EC = 1.18 EC (9.43)
9-34
-------
Table 9.2: Random Packing Costs"
Nominal _
Diameter Construction
/. , x Material
(inches)
1
1
1
2
2
3.5
3.5
304 stainless steel
ceramic
polypropylene
ceramic
polypropylene
304 stainless steel
polypropylene
Packing Type
Pall rings, Raschig rings, Bal-
last rings
Raschig rings, Berl saddles
Tri-pack®, Pall rings, Ballast
rings, Flexisaddles
Berl saddles, Raschig rings
Tri-pack®, Lanpac®, Flexir-
ing, Flexisaddle, TeUerette®
Ballast rings
Tri-pack®, Lanpac®, Ballast
rings
-. — l ' — ' —
Packing cost ($/ff'J)
< 100 ft3
70-109
33-44
14-37
13-32
3-20
30
6-14
> 100 ft:i
65-99
26-36
12-34
10-30
5-19
27
6-12
^Denotes registered trademark.
9-35
-------
9.4.2 Installation Costs
The total capital investment, TCI, is obtained by multiplying the purchased
equipment cost, PEC, by the total installation factor:
TCI = 2.20 PEC (9.44)
ThTfK,Ct°orS, ±Ci"VndUded in the total inst^tion factor are also listed
in Table 9.3.[19] The factors presented in Table 9.3 were confirmed from the
gas absorber vendor survey.
9.5 Estimating Annual Cost
The total annual cost (TAG) is the sum of the direct and indirect annual
costs.
9.5.1 Direct Annual Costs
Direct annual costs (DC) are those expenditures related to operating the
equipment, such as labor and materials. The suggested factors for each of
these costs are shown in Table 9.4. These factors were taken from Chapter 2
of this manual and were confirmed from the gas absorber vendor survey The
annual cost for each item is calculated by multiplying the number of units
used annually (i.e., hours, pounds, gallons, kWh) by the associated unit cost.
Operating labor is estimated at 1/2-hour per 8-hour shift. The super-
visory labor cost is estimated at 15 percent of the operating labor cost.
Maintenance labor is estimated at 1/2-hour per 8-hour shift. Maintenance
materials costs are assumed to equal maintenance labor costs.
Solvent costs are dependent on the total liquid throughput, the type of
solvent required, and the fraction of throughput wasted (often referred to as
blow-down). Typically, the fraction of solvent wasted varies from 0 1 percent
to 10 percent of the total solvent throughPut.[12] For acid gas systems, the
amount of solvent wasted is determined by the solids content/with bleed
off occurring when solids content reaches 10 to 15 percent to prevent salt
carry-over. 112
9-36
-------
Table 9.3: Capital Cost Factors for Gas Absorbers[19]
Cost Item Factor
Direct Costs
Purchased equipment costs
Absorber+paddng+auxiliary equipment", EC As estimated, A
Instrumentation6 0.10 A
Sales taxes 0'03 A
fteight Q.Q5 A
Purchased equipment cost, PEC B = 1.18 A
Direct installation costs
Foundations & supports
Handling & erection
Electrical
PiPin8
Insulation
Painting
Direct installation costs
Site preparation As required, SP
BuildinSs As required, Bldg.
Total Direct Costs, DC 1.85 B + SP + Bldg.
Indirect Costs (installation)
Engineering 0 10 B
Construction and field expenses 0 10 B
Contractor fees n 1 n TJ
Cl. .. u-iu a
Start-up Q 01 fi
Performance test 0 01 B
Contingencies 0 03 B
Total Indirect Costs, 1C 0.35 Q
Total Capital Investment = DC + 1C 2.20 B + SP + Bldg.
"Includes the initial quantity of packing, as well as items normally not in-
cluded with the unit supplied by vendors, such as ductwork, fan, piping, etc.
Instrumentation costs cover pH monitor and liquid level indicator in sump.
9-37
-------
Table 9.4: Suggested Annual Cost Factors for Gas Absorber Systems
Cost Item
Factor
Direct Annual Costs, DC
Operating labor0
Operator
Supervisor
Operating materials6
Solvent
Chemicals-
Wastewater disposal
Maintenance"
Labor
Material
Electricity
Fan
Pump
Indirect Annual Costs, 1C
Overhead
Administrative charges
Property tax
Insurance
Capital recovery0
Total Annual Cost
1/2 hour per shift
15% of operator
Application specific
(throughput/yr) x (waste fraction)
Based on annual consumption
(throughput/yr) x (waste fraction)
1/2 hour per shift
100% of maintenance labor
All electricity equal to:
(consumption rate) x
(hours/yr) x (unit cost)
60% of total labor and material costs
2% of Total Capital Investment
1% of Total Capital Investment
1% of Total Capital Investment
0.1315 x Total Capital Investment
DC + 1C
^These factors were confirmed by vendor contacts.
6If system does not use chemicals (e.g., caustic), this quantity is
equal to annual solvent consumption.
'Assuming a 15 year life at 10%. See Chapter 2.
9-38
-------
The total annual cost of solvent (Ca) is given by:
«n »«• / annual \ / , \
Cs = Li WF6-^ operating ( Solvent } (9 45)
hf ( hours ) \ »mi cost / ( }
where WF is the waste (make-up) fraction, and the solvent unit cost is ex-
pressed in terms of $/gal.
The cost of chemical replacement (Cc) is based on the annual consumption
of the chemical and can be calculated by:
r fibs chemical used \ / annuaj \ / chemical \
Cc = operating I cnemical
V hr / I i I unit cost j
' \ hours / \ /
where the chemical unit cost is in terms of $/lb.
Solvent disposal (Cww) costs vary depending on geographic location, type
of waste disposed of, and availability of on-site treatment. Solvent disposal
costs are calculated by:
fin «,• / annual \ / , \
Cww = lt WF™^ operating ( ,. sol-nt ) (9 47)
hr ( hours JU^POsal cost j ^47)
where the solvent disposal costs are in terms of $/gal of waste solvent.
The electricity costs associated with operating a gas absorber derive from
fan requirements to overcome the pressure drop in the column, ductwork, and
other parts of the control system, and pump requirements to recirculate the
solvent. The energy required for the fan can be calculated using Equation
y »TrO.
,., 1.17 x 10-' Gi AP
Energy/an = : (9.48)
where Energy (in kilowatts) refers to the energy needed to move a given
volumetric flow rate of air (acfm), G, is the waste gas flow rate entering
the absorber, AP is the total pressure drop through the system (inches of
H20) and e is the combined fan-motor efficiency. Values for e typically range
from 0.4 to 0.7. Likewise, the electricity required by a recycle pump can be
calculated using Equation 9.49:
_ (0-746)(2.52x 1Q-') £,(pressure)
'- (9.49)
9-39 - ' -
-------
where 0.746 is the factor used to convert horsepower to kW, pressure is
expressed in feet of water , and e is the combined pump-motor efficiency.
The cost of electricity (Ce) is then given by:
/ annual \ , .
C. = Energy . _ opg j (9.50)
where cost of electricity is expressed in units of $/KW-hr.
9.5.2 Indirect Annual Costs
Indirect annual costs (1C) include overhead, taxes, insurance, general and
administrative (G&A), and capital recovery costs. The suggested factors for
each of these items also appear in Table 9.4. Overhead is assumed to be
equal to 60 percent of the sum of operating, supervisory, and maintenance
labor, and maintenance materials. Overhead cost is discussed in Chapter 2
of this manual.
The system capital recovery cost, CRC, is based on an estimated 15-year
equipment life. (See Chapter 2 of this manual for a discussion of the capital
recovery cost.) For a 15-year life and an interest rate of 10 percent, the
capital recovery factor is 0.1315. The system capital recovery cost is then
estimated by:
CRC = 0.1315 TCI
(9.51)
G&A costs, property tax, and insurance are factored from total capital
investment, typically at 2 percent, 1 percent, and 1 percent, respectively.
9.5.3 Total Annual Cost
Total annual cost (TAG) is calculated by adding the direct annual costs and
the indirect annual costs.
TAG = DC + 1C (9.52)
9-40
-------
9.6 Example Problem #1
The example problem presented in this section shows how to apply the gas
absorber sizing and costing procedures presented in this chapter to control
a waste gas stream consisting of HC1 and air. This example problem will
use the same outlet stream parameters presented in the thermal incinerator
example problem found in Chapter 3 of this manual. The waste gas stream
entering the gas absorber is assumed to be saturated with moisture due to
being cooled in the quench chamber. The concentration of HC1 has also been
adjusted to account for the change in volume.
9.6.1 Required Information for Design
The first step in the design procedure is to specify the conditions of the gas
stream to be controlled and the desired pollutant removal efficiency. Gas and
liquid stream parameters for this example problem are listed in Table 9.5.
The quantity of HC1 can be written in terms of Ib-moles of HC1 per Ib-moles
of pollutant-free-gas (ty using the Mowing calculation:
0.001871
1 - 0.001871
Ib-moles HC1
= 0.00187
Ib-mole pollutant free gas
The solvent, a dilute aqueous solution of caustic, is assumed to have the same
physical properties as water.
9.6.2 Step 1: Determine Gas and Liquid Stream Prop-
erties
Once the properties of the waste gas stream entering the absorber are known,
the properties of the waste gas stream exiting the absorber and the liquid
streams entering and exiting the absorber need to be determined. The pol-
lutant concentration in the entering liquid (X,) is assumed to be zero. The
pollutant concentration in the exiting gas stream (Y0) is calculated using
9-41 -
-------
Table 9.5: Example Problem Data
Parameters Values
Stream Properties
Waste Gas Flow Rate Entering Absorber 21,377 scfm (22,288 acfm)
Temperature of Waste Gas Stream
Pollutant in Waste Gas
j
Concentration of HCI Entering Absorber in Waste Gas 1871 ppmv
Pollutant Removal Efficiency 99% (molar bagis)
n f rw * ^ a Water with caustic in solution
Density of Waste Gas" „ 0709 ,b/ft3
Density of Liquid[7] 62 4 ^fa
Molecular Weight of Waste Gas" 29 Ib/lb-mole
Molecular Weight of Liquid[7] 18 lb /lb.mole
Viscosity of Waste Gas" 0.044 lb/ft.hr
Viscosity of Liquid[7] 2 16 lb/ft.hr
Minimum Wetting Rate[7] ! 3 ft2 /hr
PoUutant Properties*
Diffusivity of HCI in Air 0 725 ft2 /hr
DMusivity of HCI in Water L02 x io~
Packing Properties"
Packing fype 2.inch cerkmic Raschi
Packing factor: Fp 65
Packing constant: a Z 82
Packing constant: (3 0 41
Packing constant: 7 Q'^-
Packing constant: ij> 0 0125
Packing constant: 6 0 22
Surface Area to Volume Ratio
"Reference [7], at 100°F.
'Appendix 9A.
cAppendix 9B.
Q--42
-------
Equation 9.1 and a removal efficiency of 99 percent.
/ QQ \
Y0 = 0.00187 l - = 0.0000187
The liquid flow rate entering the column is calculated from the L,/G,
ratio using Equation 9.2. Since Yj, Yot and A, are denned, the remaining un-
known, A";, is determined by consulting the equilibrium curve. A plot of the
equilibrium curve-operating line graph for an HCl-water system is presented
in Figure 9.8. The value of X; is taken at the point on the equilibrium curve
where Yj intersects the curve. The value of Yi intersects the equilibrium curve
at an X value of 0.16.
The operating line is constructed by connecting two points: (A",, Y0) and
(XZi Yi}- Tne slope of the operating line intersecting the equilibrium curve
(l,/G,)min, is:
0.00187-0.0000187
The actual L,/G, ratio is calculated using Equation 9.3. For this example,
an adjustment factor of 1.5 will be used.
pr = (0.0116)(1.5) = 0.0174
-------
0.002
0.0018
0.0002 -
f I
0.02
ace aos ai
Ib-motes HCI/lb-moles Solvent
0.14
0.16
0.18
Figure 9.8: Equilibrium Curve-Operating Line for HCI-Water SystemfT]
9-44
-------
r /'f0 « Ib-moles^ , . Ib-mc
Lmol,i = 56.8— (1 + 0) = 56.8 ^-^
\ nr / hr
>-moles
The pollutant concentration exiting the absorber in the liquid is calcu-
lated using Equation 9.10.
x = 0.00187 - 0.0000187 0.106 Ib-moles HC1
0.0174 ~ Ib-mole solvent
9.6.3 Step 2: Calculate Absorption Factor
The absorption factor is calculated from the slope of the equilibrium line and
tlle Lmol,i/Gmol,i ratio- The slope of the equilibrium curve is based on the
mole fractions of x,, *„, y;, and y0*, which are calculated from JT,-, X0, Yt",
and Y0* from Figure 9.8. From Figure 9.8, the value of Y0* in equilibrium with
the X0 value of 0.106 is 0.0001. The values of Yj* and X, are 0. The mole
fraction values are calculated from the concentration values using Equations
9.13 through 9.16.
0.106
*° = = °'096
0.0001
- °-0001
The slope of the equilibrium line from i, to !„ is calculated from Equation
y • x^r
0.0001 - 0
- m = = °-00104
Since HC1 is very soluble in water, the slope of the equilibrium curve is very
small. The absorption factor is calculated from Equation 9.11.
A n °'0174
AF = - = 17
0.00104
9.6.4 Step 3: Estimate Column Diameter
Once the inlet and outlet stream conditions are determined, the diameter of
the gas absorber may be calculated using the modified generalized pressure.
9-45
-------
drop correlation presented in Figure 9.5. The abscissa value from the graph
is calculated from Equation 9.17:
ABSCISSA = 0.0174 < = 0.000364
Since this value is outside the range of Figure 9.5, the smallest value (0.01)
will be used as a default value. The ordinate is calculated from Equation
9.19.
ORDINATE = lo-1-668-1-08^ o.oi)-o.297(log o.oi)2]
= 0.207
The superficial gas flow rate, G,frii, is calculated using Equation 9.20. For
this example calculation, 2-inch ceramic Raschig rings are selected as the
packing. The packing factors for Raschig rings are listed in Appendix 9B.
(0.207)(62.4 Ib/ft3)(0.0709 Ib/ft3)(32.2 ft/sec2)
\
0.681 lb/sec-ft2
(65)(1)(0.893)°-
2
Once Gsfr jt- is determined, the cross-sectional area of the column is cal-
culated using Equation 9.21.
4 _ (3,263 lb-mol/hr)(29 Ib/lb-mol)
(3600 sec/hr)(0.681 Ib/sec-ft2)(0.7) " 55>1 ft
The superficial liquid flow rate is determined using Equation 9.23.
, ' (56.8 lb-mol/hr)(18 Ib/lb-mol)
L»fr,i = - - - = 18'6 Ib/hr-ft2
At this point, it is necessary to determine if the liquid flow rate is sufficient
to wet the packed bed. The minimum value of L,frii is calculated using
Equation 9.24. The packing constant (a) is found in Appendix 9B.
>frJrmn = ^ ft2/hr)(62.4 Ib/ft3)(28 ft'/ff') = 2,271 lb/hr-ft2
minimum
Ttle L,fr,i value calculated using the L/G ratio is far below the
value needed to wet the packed bed. Therefore, the new value, (L • •
9-46
-------
will be used to determine the diameter of the absorber. The calculations
for this revised diameter are shown in Appendix 9C. Appendix 9C shows
that the cross-sectional area of the column is calculated to be 60 ft2, Lmoli
is 7572, and G,frii is 0.627 Ib/sec-ft2. The diameter of the column 'is then
calculated using Equation 9.22:
n (4)(60 ft2)
D = W^ 1 = 8.74 ft
V Tf
The value of X0 is then:
v 0.00187 - 0.0000187
X° = 7"572 = °-0008
Expressed in terms of mole fraction:
0.0008
1 - 0.0008
= 0.0008
The value of y0 in equilibrium with x0 cannot be estimated accurately. How-
ever, the value will approach zero, and the value of AF will be extremely
large:
„.._ 7,572
(3,263)(«0)~*°°
9.6.5 Step 4: Calculate Column Surface Area
Since x, = 0 and AF is large, Equation 9.26 will be used to calculate the
number of transfer units:
., , / 0.00187 \
Nt.. = In - 1 = 4 61
tu V0.0000187/
The height of a transfer unit is calculated from AF, HL, and H^. The
values of HG and HL are calculated from Equations 9.29 and 9.30:
ff = 3.82[(3,600)(0.7)(0.627)1"'" / Q.Q44 ' _
2,271»-'« V(0.725)(0.0709) = 2'24 ft
''22 / 2.16
- , / - - -- i r>«
2.16 ) V (0.000102)(62.4) ~
9-47
-------
The height of the transfer unit is calculated using Equation 9.28:
Htu = (2.24 ft) + —(1.06 ft) = 2.24 ft
oo
The depth of packing is calculated from Equation 9.25.
Hpack = tfft, * Htu = (4.61)(2.24 ft) = 10.3 ft
The total height of the column is calculated from Equation 9.33:
Htower = 1-40(10.3) + 1.02(8.74) + 2.81 = 26.1 ft
The surface area of the column is calculated using Equation 9.34:
S = (3.14)(8.74)(26.1 + 8.74/2) = 836 ft2
9.6.6 Step 5: Calculate Pressure Drop
The pressure drop through the column is calculated using Equation 9.35.
(0.17)(2,271) r/ft
= (n.?4)10 3.600 1(0.
0.0709
= 0.83 inches water/foot packing.
The total pressure drop (through 10.3 feet of packing) equals 8.55 inches of
water.
9.6.7 Equipment Costs
Once the system sizing parameters have been determined, the equipment
costs can be calculated. For the purpose of this example, a gas absorber
constructed of FRP will be costed using Equation 9.40.
TTC($) = 115(836) =$96, 140
The cost of 2-inch ceramic Raschig rings can be estimated from packing
cost ranges presented in Section 9.5. The volume of packing required is
calculated as:
9-48
-------
Volume of packing = (60 ft2)(10.3 ft) = 618 ft3
Using the average of the cost range for 2-inch ceramic packings, the total
cost of packing is:
Packing cost = ($20/ft3)(618 ft3) = $12,360
For this example problem, the cost of a pump will be estimated using
vendor quotes. First, the flow rate of solvent must be converted into units of
gallons per minute:
(60
V
^8.34^; \60min/
= 272 gpm
The average price for a FRP pump of this size is $16/gpm at a pressure of 60
ft water, based on the vendor survey.[12] Therefore, the cost of the recycle
pump is estimated as:
Cpvmp = (272 gpm)($16/gpm) = $4,350
For this example, the cost for a fan (FRP, backwardly-inclined centrifugal)
can be calculated using the following equation:[18]
Cfan = 57.9rf'-38
where d is the impeller (wheel) diameter of the fan expressed in inches. For
this gas flow rate and pressure drop, an impeller diameter of 33 inches is
needed. At this diameter, the cost of the fan is:
The cost of a fan motor (three-phase, carbon steel) with V-belt drive,
belt guard, and motor starter can be computed as follows:[18]
0.82T
Cmotor = 104 (hp)
As will be shown in Section 9.6.8, the electricity consumption of the fan is
32.0 kW. Converting to horsepower, we obtain a motor size of 42.6 hp. The
cost of the fan motor is:
C motor = 104(42.6)"-821 = $2,260
9-49
-------
The total auxiliary equipment cost is:
$4,350 + $7,210 -I- $2,260 = $13,820
The total equipment cost is the sum of the absorber cost, the packing
cost, and the auxiliary equipment cost:
EC = 96,140 + 12,360 + 13,820 = $122,320
The purchased equipment cost including instrumentation, controls, taxes,
and freight is estimated using Equation 9.43:
PEC = 1.18(122,320) = $144,340
The total capital investment is calculated using Equation 9.44:
TCI = 2.20(144,340) = $317,550 w $318,000
9.6.8 Total Annual Cost
Table 9.6 summarizes the estimated annual costs using the suggested factors
and unit costs for the example problem.
Direct annual costs for gas absorber systems include labor, materials,
utilities, and wastewater disposal. Labor costs are based on 8,000 hr/year of
operation. Supervisory labor is computed at 15 percent of operating labor,
and operating and maintenance labor are each based on 1/2 hr per 8-hr shift.
The electricity required to run the fan is calculated using Equation 9.48
and assuming a combined fan-motor efficiency of 70 percent:
,, (1.17 x 10-')(22,288)(8.55)
Energy/an = ' -1 L = 32.0 kW
The energy required for the liquid pump is calculated using Equation 9.49.
The capital cost of the pump was calculated using data supplied by vendors
9-50
-------
Table 9.6: Annual Costs for Packed Tower Absorber
Example Problem
Item
Calculations
Cost
Direct Annual Coats, DC
Operating Labor
Operator
Supervisor
Operating materials
Solvent (water)
Caustic Replacement
Wastewater disposal
Maintenance
Labor
Material
Electricity
Total DC
Indirect Annual Costs, 1C
Overhead
Administrative charges
Property tax
Insurance
Capital recovery"
Total IC
fl* ><
15% of operator = 0.15 x 7,820
7.16 gpm x
3.06 Ib-mole 62jb 8,000 hr ton
hr Ib-mole yr x 2000 IB
0776 x Ton"
7.16 gpm x
100% of maintenance labor
36.4 kW x $iOOOhr x 80.0461
60% of total labor and maintenance material:
= 0.6(7,820 + 1,170 + 8,610 + 8,610)
2% of Total Capital Investment = 0.02(8317,550)
1% of Total Capital Investment = 0.01(8317,550)
1% of Total Capital Investment = 0.01(8317,550)
0.1315 x $317,550
Total Annual Cost (rounded)
$7,820
1,170
690
13i060
8,610
$352,940
15,730
6,350
3,180
3,180
41,760
8423 000
"The capital recovery cost factor, CRF, is a function of the absorber equipment life and
the opportunity cost of the capital (i.e., interest rate). For this example, assume a 15-year
equipment life and a 10% interest rate.
9-51
-------
for a pump operating at a pressure of 60 feet of water. Assuming a pressure
of 60 ft of water and a combined pump-motor efficiency of 70 percent:
Energy _ (0-746) (2.52 x 1Q-»)(272)(60)(1)
^pump — T-rr - — = 4.4 kW
The total energy required to operate the auxiliary equipment is approxi-
mately 36.4 kW. The cost of electricity, Ce, is calculated using Equation 9.50
and with the cost per kWh shown in Table 9.6.
Ce = (36.4 kW)(8,000 h/yr)($0.0461/kWh) = $13,420/yr
The costs of solvent (water), wastewater disposal, and caustic are all
dependent on the total system throughput and the fraction of solvent dis-
charged as waste. A certain amount of solvent will be wasted and replaced
by a fresh solution of water and caustic in order to maintain the system's
pH and solids content at acceptable levels. Based on the vendor survey, a
maximum solids content of 10 percent by weight will be the design basis for
this example problem.[12] The Mowing calculations illustrate the procedure
used to calculate how much water and caustic are needed, and how much
solvent must be bled off to maintain system operability.
From previous calculations, Lmol>i = 7,572 Ib-moles/lir. The mass flow
rate is calculated as:
T (7 *7o lb-mole\ / ib \ lb
L= 7572— - - - -
m-mole HCIW lb \ lb HCI
Gma33,HCl = 6.12 - - - 36.5 - - - = 223.4 ^-^
V nr ) \ lb-mole/ hr
9-52
-------
For this example problem, the caustic is assumed to be Na20, with one
mole of caustic required for neutralizing 2 moles of HCL. Therefore, 3.06
Ib-moles/hr of caustic are required.
The unit cost of a 76 percent solution of Na20 is given in Table 9.6. The
annual cost is calculated from:
Cc =
__
hr J\ lb-mole,/V 3* A2,000 \b) \0,7Qj \ ton I
= $299, 560 yr
Mass of the salt formed in this chemical reaction, NaCl, is calculated as:
MaSS = ( 223 jlb'HC1>l ( lb-mole "\ /Hb-moleNaCl.\ ( 58.5 Ib NaCl \
NaC1 V ' hr A36.51bHClA Ib-mole HC1 Alb-mole NaCl J
hr
If the maximum concentration of NaCl in the wastewater (ww) is assumed
to be 10 weight percent, the wastewater volume flow rate is calculated as:
Wastewater = / Ib NaCA / 1 Ib ww \ / gal ww \ 1 hi
fl°Wrate V ' hr ; 1,0.1 Ib NaCiy 1,8.34 Ib ww J
= 7.16 gpm
where 8.34 is the density of the wastewater.
The cost of wastewater disposal is:1
The cost of solvent (water) is:
r C7i« \ n r .
C. = (T.16 gpn.) -j_ 8,000- — = »690/yr
'Because the wastewater stream contains only NaCl, it probably will not require pre-
treatment before discharge to a municipal wastewater treatment facility Therefore the
^St8nW/inenndiSn°Sal T* C°St Sh°Wn h"e " just a Sewer Usa8e rate' This ^^
n North r gr ° "tin ^r"^6 °f thC mteS Chatged by the SCVen la^est municipalities
m Worth Carolma.[20] These rates range from approximately S2 to $6/1000 gal This wide
range is indicative of the major differences among sewer rates throughout the country.
9-53
-------
Indirect annual costs include overhead, administrative charges, property
tax, insurance, and capital recovery. Total annual cost is estimated using
Equation 9.52. For this example case, the total annual cost is estimated to
be $423,000 per year (Table 9.6).
9.7 Example Problem #2
In this example problem the diameter of a gas absorber will be estimated
by defining a pressure drop. A pressure drop of 1 inch of water per foot
of packing will be used in this example calculation. Equation 9.38 will be
used to calculate the ordinate value relating to an abscissa value. If the
Lmol,i/Gmol,i rati° ls known, the abscissa can be calculated directly. The
ordinate value is then:
ORDINATE = exp [-4.0950 - 1.0012 ln(0.0496) - 0.1587(ln 0.0496)2+
0.0080(ln 0.0496)3 + 0.0032(ln 0.0496)']
= 0.084
The value of G^ is calculated using Equation 9.39
Gsfr,i =
\
(62.4 - 0.0709)(0.07Q9)(32.2)(O.Q84)
65(0.893)°
.1
= 0.43 Ib/ft2-sec
The remaining calculations are the same as in Section 9.3.4, except the flood-
ing factor is not used in the equations.
9.8 Acknowledgements
The authors gratefully acknowledge the Mowing companies for contributing
data to this chapter:
• Air Plastics, Inc. (Cincinnati, OH)
• Airpol, Inc. (Teterboro, NJ)
-9-54
-------
• Anderson 2000, Inc. (Peachtree City, GA)
• Calvert Environmental (San Diego, CA)
• Ceilcote Air Pollution Control (Berea, OH)
• Croll-Reynolds Company, Inc. (Westfield, NJ)
• Ecolotreat Process Equipment (Toledo, OH)
• Glitsch, Inc. (Dallas, TX)
• Interel Corporation (Englewood, CO)
• Jaeger Products, Inc. (Spring, TX)
• Koch Engineering Co., Inc. (Wichita, KS)
• Lantec Products, Inc. (Agoura Hills, CA)
• Midwest Air Products Co., Inc. (Owosso, MI)
• Monroe Environmental Corp., (Monroe, MI)
• Norton Chemical Process Products (Akron, OH)
9-55
-------
-------
Appendix 9A
Properties of Pollutants
9-56
-------
-------
Table 9.7: Physical Properties of Common Pollutants0
Pollutant
Ammonia
Methanol
Ethyl Alcohol
Propyl Alcohol
Butyl Alcohol
Acetic Acid
Hydrogen Chloride
Hydrogen Bromide
Hydrogen Fluoride
Molecular
Weight
Gblb )
17
32
46
60
74
60
36
36
20
Diffusivity in
Air
at 25° C
(cm2/sec)
0.236
0.159
0.119
0.100
0.09
0.133
0.187
0.129
0.753
Diffusivity in
Water
at 20° C
(cm2/sec)xl05
1.76
1.28
1.00
0.87
0.77
0.88
2.64
1.93
3.33
"Diffusivity data taken from Reference [7, 21].
9-57
-------
-------
Appendix 9B
Packing Characteristics
-9-58
-------
-------
Table 9.8: Packing Factors for Various Packings[3, 7, 10, 13]
Packing
Type
Raschig rings
Raschig rings
Pall rings
Pall rings
Berl saddles
Intalox saddles
Tri-Packs®
Construction
Material
ceramic
metal
metal
polypropylene
ceramic
ceramic
plastic
Nominal
Diameter
(inches)
1/2
5/8
3/4
1
1 1/2
2
3
1/2
5/8
3/4
1
1 1/2
2
3
5/8
1
1 1/2
2
3 1/2
5/8
1
1 1/2
2
1/2
3/4
1
1 1/2
2
1/2
3/4
1
FP
640
380
255
160
95
65
37
410
290
230
137
83
57
32
70
48
28
20
16
97
52
32
25
240
170
110
65
45
200
145
98
1 1/2 52
2
3
2
3 1/2
40
22
16
12
a
111
100
80
58
38
28
118
72
57
41
31
21
131
66
48
36
110
63
39
31
142
82
76
44
32
190
102
78
60
36
48
38
9-59
-------
Table 9.9: Packing Constants Used to Estimate #G[1, 3, 7, 13]
Packing
Type
Raschig Rings
Berl Saddles
Partition Rings
~~ M\
LanPac®
Tri-Packs®
Size
(inches)
—
3/8
1
1
1 1/2
1 1/2
2
1/2
1/2
1
1 1/2
3
2.3
2
3 1/2
— ^ ^ — —
Packing Constants
« | 0
2.32
7.00
6.41
1.73
2.58
3.82
32.4
0.81
1.97
5.05
640.
7.6
1.4
1.7
0.45
0.39
0.32
0.38
0.38
0.41
0.30
0.30
0.36
0.32
0.58
0.33
0.33
0.33
7
0.47
0.58
0.51
0.66
0.40
0.45
0.74
0.24
0.40
0.45
1.06
-0.48
0.40
0.45
Applicable Range"
<**
200-500
200-800
200-600
200-700
200-700
200-800
200-700
200-700
200-800
200-1,000
150-900
400-3,000
100-900
100-2,000
L,fr
500-1,500
400-500
500-4,500
500-1,500
1,500-4,500
500-4,500
500-1,500
1,500-4,500
400-4,500
400-4,500
3,000-10,000
500-8,000
500-10,000-
500-10,000
9-60
-------
Table 9.10: Packing Constants Used to Estimate HL[1, 3, 13]
Packing
Type
Raschig Rings
fieri Saddles
Partition Rings
LanPac®
Tri-packs®
Size
(inches)
3/8
1
1 1/2
21/2
2
1/2
1
1 1/2
3
2.3
3.5
2
3 1/2
Packing Constants
<£
0.00182
0.00357
0.0100
0.0111
0.0125
0.00666
0.00588
0.00625
0.0625
0.0039
0.0042
0.0031
0.0040
6
0.46
0.35
0.22
0.22
0.22
0.28
0.28
0.28
0.09
0.33
0.33
0.33
0.33
Applicable Range
Tn
400-15,000
400-15,000
400-15,000
400-15,000
400-15,000
400-15,000
400-15,000
400-15,000
3,000-14,000
500-8,000
500-8,000
500-10,000
500-10,000
0 Units of lb/hr-ft2
9-61
-------
Table 0.11: Packing Constants Used to Estimate Pressure Drop[l, 7, 13]
Packing
Type
Raschig rings
Raschig rings
Pall rings
Berl saddles
Intalox saddles
Construction
Material
ceramic
metal
metal
ceramic
ceramic
Nominal
Diameter
(inches)
1/2
3/4
1
1 1/4
11/2
2
5/8
1
1 1/2
2
5/8
1
1 1/2
2
1/2
3/4
1
1 1/2
1/2
3/4
1
1 1/2
c J
3.1
1.34
0.97
0.57
0.39
0.24
1.2
0.42
0.29
0.23
0.43
0.15
0.08
0.06
1.2
0.62
0.39
0.21
0.82
0.28
0.31
0.14
0.41
0.26
0.25
0.23
0.23
0.17
0.28
0.21
0.20
0.135
0.17
0.16
0.15
0.12
0.21
0.17
0.17
0.13
0.20
0.16
0.16
0.14
9-62
C
-------
Appendix 9C
Minimum Wetting Rate
Analysis
As explained m the design procedures, the liquid flow rate entering the col-
umn must be high enough to effectively wet the packing. If the liquid flow
rate as determined theoretically in Equation 9.23, is lower than the flow rate
dictated by the minimum wetting rate, calculated in Equation 9.24, then the
packing will not be wetted sufficiently to ensure mass, transfer between the
gas and liquid phases. The minimum liquid flow rate should then be used as
a default value. The superficial gas flow rate, (^ , and cross-sectional area
of the column must then be recalculated to account for the increased liquid
flow rate. The approach necessary to recalculate these variables is explained
m Section 9C.1 of this Appendix. The calculation of these variables using
the results from Example Problem #1 are presented in Section 9C 2 of this
Appendix.
9C.1 Overview of the Approach
1. The value of Lmolj. must be recalculated from the value of (L
using the equation: V '
L ,.-
moi'1
MWL
9-63
-------
The value of A (the cross-sectional area of the absorber column) is the
only unknown in the equation.
2. The ABSCISSA value is calculated in terms of A by substituting the
new Lmoiii into Equation 9.17.
3. The value of Gafr)t- is recalculated by rearranging Equation 9.21, with
A as the only unknown.
4. The ORDINATE value is calculated in terms of A from the new Gtfr i
using Equation 9.18.
5. An iterative process is used to determine A, ABSCISSA, and ORDI-
NATE. Values of A are chosen and the ABSCISSA and ORDINATE
values are calculated. The ORDINATE value corresponding to the AB-
SCISSA value is determined from Figure 9.5 (or Equation 9.19), and
this value is compared to the ORDINATE value calculated using Equa-
tion 9.18. This process is continued until both ORDINATE values are
equal.
9C.2 Example Problem Calculation
Step 1: The first step is to recalculate the liquid flow rate. The liquid molar
flow rate may be calculated using Equation 9.23.
Lmolii = (
= (126.2 Ib-mole/hr-ft2)/!
Step 2: The abscissa value from Figure 9.5, and presented in Equation 9.17, is
calculated as:
ABSCISSA = (126-2 Ib-mole/hr-ft2)>i /18\ /Q.0709
3,263 Ib-mole/hr V29/ V 62.4
= 8.09 x W-4A (9.53)
Step 3: The value of Gsfpii is then recalculated in terms of the cross-sectional
area of the column.
G = (3,263 lb-mole/hr)(29 Ib/lb-mole) _ 3_7.6
3fr (3600 s ~~
9-64
-------
Step 4: The ordinate value from Figure 9.5, and presented in Equation 9.18 is
calculated as: '
ORDINATE =
(62.4)(0.0709)(32.2)
631
A2
(9.54)
Step 5: At this point the simplest solution is an iterative approach. Choose
a value for A, calculate the ABSCISSA value using Equation 9.53,
and find the corresponding ORDINATE value off the flooding curve in
Figure 9.5 (or use Equation 9.19 to calculate the ORDINATE value)
Compare the calculated ORDINATE value from Equation 9.54 to the
value obtained from the graph or from Equation 9.19. By continuing
this process until the ORDINATE values converge the value of 4 is
determined to be 60 ft2. The Mowing table illustrates the intermediate
steps in the calculational process.
Assumed
Value
of A
65
62
60
ABSCISSA
Calculated
From Eqn. 9.53
0.0526
0.0502
0.0485
ORDINATE
Calculated
From Eqn. 9.19
0.1714
0.1740
0.1757
ORDINATE
Calculated
From Eqn. 9.54
0.1493
0.1642 •
0.1752
The value of
is then:
The liquid molar flow rate is:
Lmol,i = (126.2)(60) = 7,572 Ib-mole/hr
The diameter and height of the column using the results of this calculation
are presented in Example Problem #1.
9-65
-------
-------
References
[1] Control Technologies for Hazardous Air Pollutants, Office of Research
and Development, U.S. Environmental Protection Agency, Research Tri-
angle Park, North Carolina, Publication No. EPA 625/6-91-014.
[2] Mclnnes, R., K. Jameson, and D. Austin, "Scrubbing Toxic Inorganics",
Chemical Engineering, September 1990, pp. 116-121.
[3] Letter from Jose L. Bravo of Jaeger Products, Inc., to William M.
Vatavuk, U.S. Environmental Protection Agency, June 8, 1992.
[4] Treybal, Robert E., Mass Transfer Operations (Third edition), McGraw-
Hill Book Company, New York, 1980.
[5] Letter from Jack D. Brady of Anderson 2000, Inc., to William M.
Vatavuk, U.S. Environmental Protection Agency, June 9, 1992.
[6] Letter from S. Raymond Woll of Air Products, Inc., to William M.
Vatavuk, U.S. Environmental Protection Agency, June 25, 1992.
[7] Perry, R.H. and C.H. Chilton, Eds., Chemical Engineers' Handbook
(Sixth edition), McGraw-Hill Book Company, New York, 1984.
[8] Crowe, Charles R., and D. Cooper, "Brick/Membrane Linings Pass the
Acid Test", Chemical Engineering, July 1988, pp. 83-86.
[9] Harrison, Mark E., and John J. France, "Distillation Column Trou-
bleshooting, Part 2: Packed Columns", Chemical Engineering, April
T10I Coker, A.K., "Understanding the Basics of Packed-Column Design".
Chemical Engineering Progress, November 1991,,pp. 93-99.
[11] Telephone conversation between Roy Oommen, Radian Corporation and
Gerald Nealon, Norton Process Equipment, April 4, 1992.
9-66
-------
[12] Gas absorber questionnaire responses from nine gas absorber vendors to
Radian Corporation, August-December, 1991.
[13] Buonicore, A. J., and L. Theodore, Industrial Control Equipment for
Gaseous Pollutants, Volume I, CRC Press, Inc., Cleveland, Ohio, 1975.
[14] Strigle, Ralph F., Random Packings and Packed Towers, Design Appli-
cations, Gulf Publishing Company, Houston, Texas, 1987.
[15] Questionnaire response from K. C. Lang of Lantec Products, Inc. to
R.V. Oommen, Radian Corporation, August 23, 1991.
[16] Memorandum from Vatavuk, W.M. of U.S. Environmental Protection
Agency to Martha Smith, U.S. EPA, March 27, 1992.
[17] Packing vendor questionnaire responses from seven packing vendors to
Radian Corporation, August, 1991-January, 1992.
[18] Vatavuk, W.M., "Pricing Equipment for Air-Pollution Control", Chem-
ical Engineering, May 1990, pp. 126-130.
[19] Vatavuk, W.M., and R.B. Neveril, "Estimating Costs of Pollution Con-
trol Systems, Part II: Factors for Estimating Capital and Operating
Costs", Chemical Engineering, November 3, 1980, pp. 157-162.
[20] Telephone conversation between William M. Vatavuk, U.S. Environmen-
tal Protection Agency, and Cindy Kling, City of Raleigh, N.C., July 16,
[21] "Air Pollution Engineering Manual" (AP-40), (Second Edition), Daniel-
son, John A., Los Angeles County Air Pollution Control District CA
May 1973.
9-67
-------
Chapter 10
HOODS, DUCTWORK, and STACKS
William M. Vatavuk
Standards Development Branch, OAQPS
U.S. Environmental Protection Agency
Research Triangle Park, NC 27711
March 1994
-------
Contents
. . 10-3
10.1 Introduction ................
. ,_ . . . 10-4
10.2 Equipment Description ............. 1Q_4
10.2.1 Hoods .................... 10_4
10.2.1.1 Types of Hoods ............ io_Q
10.2.2 Ductwork ................... in in
10.2.2.1 Ductwork Components .......... tn 1-3
J-U - -L-J
10.2.3 Stacks ................
. . 10-14
10.3 Design Procedures ............... 10-15
10.3.1 Design Fundamentals .............
10.3.1.1 The Bernoulli Equation -,-••;•••• ^ j
10 3 1 2 Pressure: Static, Velocity, and Total 10-18
10 ".a! 1.3 Temperature and Pressure Adjustments . 10-21
10.3.2 Hood Design Procedure '
10 3.2.1 Hood Design Factors
n
"
.
10.3.2.2 Hood Sizing Procedure n 29
10.3.3 Ductwork Design Procedure . . . . ...... ^-
10 3 3 1 Two Ductwork Design Approaches . . - - ^"^
10 '3 3.2 Ductwork Design Parameters ...... 10- ^
10.3.3.3 Ductwork Pressure Drop ........ tn" 37
10 3.4 Stack Design Procedures ........... in\«
10 3.4.1 Calculating Stack Diameter ...... iu-J«
10.3.4.2 Calculating Stack Height ....... 10- J«
10.3.4.3 Calculating Stack Draft ........ lu-^u
10.4 Estimating Total Capital Investment ......... 10-41
10.4.1 Equipment Costs ...............
10.4.1.1 Hood Costs .............. ^
10.4.1.2 Ductwork Costs ............ ^
10 4 1.3 Stack Costs .............. ^ °
10.4.2 Taxes, Freight, and Instrumentation Costs . . 10-53
10.4.3 Purchased Equipment Cost ........... 10 54
10.4.4 Installation Costs ..............
10 5 Estimating Total Annual Cost ............ lo'ss
10.5.1 Direct Annual Costs ............. 10 56
10.5.2 Indirect Annual Costs ............ 10-56
10.5.3 Total Annual Cost ..............
. . . 10-57
10.6 Acknowledgements ...............
. . . 10-58
References ..................
10-2
-------
-------
10.1 Introduction
Most control devices are located some distance from the
emission sources they control. This separation may be needed for
several reasons. For one thing, there may not be enough room to
install the control device close to the source. Or, the device
may collect emissions from several sources located throughout the
facility and, hence, must be sited at some convenient,
equidistant location. Or, it may be that required utility
connections for the control device are only available at some
remote site. Regardless of the reason, the waste gas stream must
be conveyed from the source to the control device and from there
to a stack before it can be released to the atmosphere.
The kinds of equipment needed to convey the waste gas are
the same for most kinds of control devices. These are: (1)
hoods, (2) ductwork, (3) stacks, and (4) fans. Together, these
items comprise a ventilation system. A hood is used to capture
the emissions at the source; ductwork, to convey them to the
control device; a stack, to disperse them after they leave the
device- and a fan, to provide the energy for moving them through
the control system. This chapter covers the first three kinds of
equipment. However, because they constitute such a broad and
complex subject, fans will be dealt with in a future Manual
chapter Also, the kinds of stacks covered are short stacks
(100-120 feet high or less). Typically, these are included with
packaged control systems or added to them. So-called "tall
stacks" ("chimneys"), used at power plants or other sources where
the exhaust gases must be dispersed over great distances, will -
not be discussed in this chapter.
This chapter presents all the information one would need to
develop study (+ 30%-accurate) cost estimates for hoods,
ductwork, and stacks. Accordingly, the following sections
include- (1) descriptions of the types of equipment used in air
pollution control ventilation systems, (2) procedures for sizing
(designing) this equipment, and (3) methodologies and data for
estimating their capital and annual costs. Also, sprinkled
throughout the chapter are several illustrations (example
problems) that show the reader how to apply the various sizing
and costing methodologies.
10-3
-------
10.2 Equipment Description
In this section, the kinds of hoods, ductwork, and stacks
used in air pollution control systems are described, each in a
separate subsection. These descriptions have been based on _
information obtained from standard ventilation and air pollution
control references, journal articles, and equipment vendors.
10.2.1 Hoods
Of the several components of an air pollution control
system, the capture device is the most important. This should be
self-evident for if emissions are not efficiently captured at
the source they cannot be conveyed to and removed by a control
device There are two general categories of capture devices: (1)
direct exhaust connections (DEC) and (2) hoods. As the name
implies, a DEC is a section of duct (typically an^elbow) into
which the emissions directly flow. These connections often are
used when the emission source is itself a duct or vent such as a
process vent in a chemical manufacturing plant or petroleum
refinery. (See discussion below on "Ductwork".)
Hoods comprise a much broader category than DECs. They are
used to capture particulates, gases, and/or mists emitted from a
variety of sources, such as basic oxygen steelmaking furnaces,
welding operations, and electroplating tanks. The hooded
processes are generally categorized as either "hot" or "cold" a
delineation that, in turn, influences hood selection, placement.-
and design.
The source conditions also influence the materials from _
which a hood is fabricated. Mild (carbon) steel is the material
of choice for those applications where the emission stream is
noncorrosive and of moderate temperature. However, where
corrosive substances (e.g., acid gases) are present in high
enough concentrations, stainless steels or plastics (e.g.,
l?berglass-reinforced plastic, orFRP) are_needed As most hoods
are custom-designed and built, the vendor involved would
determine which material would be optimal for a given
application.
10.2.1.1 Types of Hoods
Although the names of certain hoods vary, depending on which
ventilation source one consults, there is general agreement as to
how they are classified. There are four types of hoods: (1)
enclosures, (2) booths, (3) captor (capturej hoods, and (4)
receptor (receiving) hoods.1'2
10-4
-------
Enclosures are of two types: (1) those that are Completely
closed to the outside environment and (2) those that have
openings for material input/output. The Jirst. W6^^yby
when handling radioactive materials which ™«t be ^andled by
remote manipulators. They are also dust- and gas-tigh^ These
kinds of enclosures are rarely used in air pollution control.
Total enclosures, the second type, have applications in .
c^PVPral areas such as the control of emissions from electric arc
furnaces and from screening and bin filling operations. They are
equipped with small wall openings (natural draft
ooeninqs— "NDO's") that allow for material to be moved in or out
and foTvenSlation. However, the area of these openings must be
small compared with the total area of the enclosure walls
(typically, 5% or less).
Another application of total enclosures is in the
measurement ofP?he capture efficiency of VOC (volatile organic
compound) control devices. Capture efficiency a?that fraction
of all VOC's generated at, and released by, an affected facility
?hat is directed to the control device. In this application, a
total enclosure is a temporary structure that completely
sSrrounSs an emitting process so that all VOC emissions are
captured for discharge through ducts or stacks The air flow
through the total enclosure must be high enough to keep the
concentration of the VOC mixture inside the enclosure within both
tSe Sectional Safety and Health Administration (OSHA) health
requirement limits and the vapor explosive limits. (The latter
are typically set at 25% of the lower explosive limit (LED for
?he VOC mSture in question.) In addition, the overall face
velocity of air flowing through the enclosure must be at least -
200 ft/min.3
The surfaces of temporary total enclosures are usually
constructed either of plastic film or of such rigid materials as
insulSiofpJnels or pfywood. Plastic f ilm. of f ere the ad vantages
of beina lightweight, transparent, inexpensive, and easy to work
with However, i? is flimsy, flammable, and has a relatively low
melting point. In addition, the plastic must be hung on a
framework of wood, plastic piping, or scaffolding.
Although rigid materials are more expensive and less
workable than plastic, they are more durable and can_withstand
larger pressure differentials between the enclosure interior and
exterio? Total enclosure design specifications (which have been
incorporated into several EPA emission standards) are contained
in the EPA report, The Measurement Solution: Usmg a Temporary
Total Enclosure for Capture Testing.4
Booths are like enclosures~ in that they surround the
emission source, except for a wall (or portion thereof) that is
omitted to allow access by operators and equipment. Like
10-5
-------
enclosures, booths must be large enough to prevent particulates
from impinging on the inner walls. They are used with such
operations (and emission sources) as spray painting and portable
grinding, polishing, and buffing operations.
Captor Hoods: Unlike enclosures and booths, captor hoods
(also termed active or external hoods)_ do not enclose the source
at all Consisting of one to three sides, they are located at a
distance from the source and draw the emissions into them via
fans Captor hoods are further classified as side-draft/back-
draft slot, downdraft, and high-velocity, low-volume (HVLV)
hoods. A side-draft/back-draft hood is typically located to the
side/behind of an emission source, but as close to it as
possible as air velocities decrease inversely (arid sharply) with
distance. Examples of these include snorkel-type welding hoods
and side shake-out hoods.
A slot hood operates in a manner similar to a side-
draft/back-draft. However, the inlet opening (face) is_much
smaller, being long and narrow. Moreover, a slot hood is
situated at the periphery of an emission source, such as a
narrow, open tank. This type of hood is also employed with bench
welding operations.
While slot and side-draft/back-draft hoods are located
beside/behind a source, a downdraft hood is situated immediately
beneath it It draws pollutant-laden air down through the source
and thence, to a control device. Applications ot down-draft
hoods include foundry shake-out and bench soldering and torch
cutting operations.
HVLV hoods are characterized by the use of extremely high
velocities (capture velocities) to collect contaminants at the
source and by the optimal distribution of those velocities
across'the hood face. To maintain a low volumetric flow rate,
these hoods are located as close to the source as possible, so as
to minimize air entrainment.
Receptor hoods: The last category is receptor hoods (a.k a.
passive or canopy hoods). A receptor hood typically is located
above or beside a source, to collect the emissions, which are
given momentum by the source. For example, a canopy hood might
be situated directly above an open tank containing a hot liquid
(a buoyant source) . With entrained air, vapors emitted from the
liquid would rise into the hood. Here, the canopy hood would
function as a passive collector, as the rising gases would be
drawn into the hood via natural draft. (See Figure 10.1.)
Receptor hoods are-also used with nonbuoyant sources,
sources from which emissions do not rise. However, the emissions
can be "thrown off" from a process, such as a swing grinder. The
initial velocity of the emissions typically is high enough to
10-6
-------
Figure 10.1 Typical Canopy Hood Installation
Source: tank or process
10-7
-------
convey them into a receiving hood.5
10.2.2 Ductwork
Once the emission stream is captured by either a hood or a
direct exhaust connection, it is conveyed to the control device
via ductwork. The term "ductwork" denotes all of the equipment
between the capture device and the control device. This
includes: (1) straight duct; (2) fittings, such as elbows and
tees; (3) flow control devices (e.g., dampers); and (4) duct
supports. These components are described in Section 10.2.2.1.)
In air pollution control systems, the fan is usually located
immediately before or after the control device. Consequently,
most of the ductwork typically is under a negative static
pressure, varying from a few inches to approximately 20 inches of
water column. These pressure conditions dictate the type of duct
used as well as such design parameters as the wall thickness
(gauge). For instance, welded duct is preferable to spiral-wound
duct in vacuum applications.6
Ductwork is fabricated from either metal or plastic, the
choice of material being dictated by the characteristics of the
waste gas stream, structural considerations, purchase and
installation costs, aesthetics, and other factors. Metals used
include carbon steel- (bare or galvanized), stainless steel, and
aluminum. The most commonly used plastics are PVC (poly-vinyl
chloride) and FRP (fiberglass-reinforced plastic), although
polypropylene (PP) and linear polyethylene (LPE) also have been
applied. However, one serious drawback to PP and LPE is that
both are combustible.
PVC and other plastic ductwork are resistant to a variety of
corrosive substances, from aqua regia to 95% sulfuric acid. But
plastic ductwork cannot tolerate environmental temperatures above
150°F.8 Metal ductwork can handle temperatures up to
approximately 1000°F, but only certain alloys can tolerate
corrosive streams.
In terms of construction, ductwork can be either rigid or
flexible As the name implies, rigid ductwork, whether metal or
plastic, has a fixed shape. Conversely, flexible ductwork can be
bent to accomodate situations where space is limited or where the
layout is so convoluted that rigid fittings cannot meet
construction requirements. Usually circular in cross-sectional
shape, flexible duct can be fabricated from metals or plastic and
can be either insulated or uninsulated.
Rigid ductwork is fabricated into circular, flat oval, or
square/rectangular cross-sectional shapes. Of these, circular
duct is most commonly used in air pollution control systems.
-------
Although square/rectangular duct is advantageous to use when
space is limited, round duct offers several advantages. It
resists collapsing, provides better transport conditions, and
uses less metal than square/rectangular or flat oval shapes of
equivalent cross-sectional area.9 Unless otherwise noted, the
following discussion will pertain to rigid, circular duct, as
this is the type most commonly used in air pollution control.
Rigid metal circular duct is further classified according_to
method of fabrication. Longitudinal seam duct is made by bending
sheet metal into a circular shape over a mandrel, and butt-
welding the two ends together. Spiral seam duct is constructed
from a long strip of sheet metal, the edges of which are joined
by an interlocking helical seam that runs the length of the duct.
This seam is either raised or flush to the duct wall surface.
Fabrication method and cross-sectional shape are not the
only considerations in designing ductwork, however. One must
also specify the diameter; wall thickness; type, number, and
location of fittings, controllers, and supports; and other
parameters Consequently, most ductwork components are custom-
designed and fabricated, so as to optimally serve the control
device. Some vendors offer prefabricated components, but these
are usually common fittings (e.g., 90° elbows) that are available
only in standard sizes (e.g., 3- to 12-inch diameter) • .
If either the gas stream temperature or moisture content is
excessive the ductwork may need to be insulated. Insulation
inhibits heat loss/gain, saving energy (and money), on the one
hand and prevents condensation, on the other. Insulation also.
protects personnel who might touch the ductwork from sustaining
burns There are two ways to insulate ductwork. The first is to
install insulation on the outer surface of the ductwork and cover
it with a vapor barrier of plastic or metal foil. The type and
thickness of insulation used will depend on several heat
transfer-related parameters. For instance, one vendor states
that 4 inches of mineral wool insulation is adequate for
maintaining a surface ("skin") temperature of 140°F (the OSHA
workplace limit) or lower, provided that the exhaust gas
temperature does not exceed 600°F.12
The second way to insulate ductwork is by using double-wall,
insulated duct and fittings. Double-wall ductwork serves to
reduce both heat loss and noise. One vendor constructs it from a
solid sheet metal outer pressure shell and a sheet metal inner
liner with a layer of fiberglass insulation sandwiched between.
The insulation layer is typically 1-inch, although 2- and 3-inch
thicknesses are available for more extreme applications. The
thermal conductivities of these thicknesses are 0.27, 0.13, and
0.09 Btu/hr-ft2-°F, respectively.13
10-9
-------
10.2.2.1 Ductwork Components
As discussed above, a ductwork system consists of straight
duct, fittings, flow control devices, and supports^ Straight
stresses arise.
The most commonly used fittings are e^°"* (:f^"> Q0 Th|fe
e mo Q0
this directional change occurs, ^'f^^^^fe^ional
centerlme radius (Rd) is 1.5 x c ne e elbows in which the
in stanird el.ows, „, -
> 2Dc-
Tees are used when two or more gas streams must be
f f are u streams converge at a 90
^osse^are'also SsedCtc. connfct^ct .ranches. Here, the «o
branches intersect each other at a nght angle.
,«
the diameter decreases wholly on one side of the fitting.
To control the volumetric flowrate through ventilation
10 coriuiu-L unc r^mT-^-rc! a TP usuallv delineated
"
_1_ \-4 J. 1 k_ W -i- j. j. _»- ^-j — — / _-_- ^ . ,
macic;. In single blade dampers, a
to a rod, one end of which protrudes
• commonly used type of single blade
outside the duct In tne ^sed ^ control the gas flow
damper (.butterfly type , ^1J-° ^^^^^ pnilv closed the damper
by rotating the plate in the damper^ ^^eft°^' fully opL,
theefIceSirpa?a?lel to tSe gas^low lines. Several-single blade
"control" dampers are depicted in Figure 10.2.
10-10
-------
Figure 10.2 Selected Circular Ductwork Componentsf
LONGITUDINAL
SEAM DUCT
(Fully welded
lonaitudmal seam)
DIMENSIONS.
90" maximum
GORED ELBOW
DIMENSIONS:
STRAIGHT TEE
STRAIGHT 90° CROSS
DIMENSIONS:
V - C - 2
Maximum C = A
DIMENSIONS:
V - C » 2
Maximum Cor O * A
HEAVY-DUTY
CONTROL DAMPER
CONCENTRIC
REDUCER
ECCENTRIC REDUCER
f Reference: "Single-Wall Round and Flat Oval Duct and
Fittings." In: Sheet Metal Division Catalog. Groveport, OH:
United McGill Corporation. 1990.
10-11
-------
With blast gate dampers, a second type, the flow is
controlled by sliding the damper blade in and out of the duct.
Blast gates are often used to control the flow of air streams
containing suspended soiids, such as in pneumatic conveyors. In
tSesJ Respects butterfly dampers and blast gates are analogous,
resnectlvSy to the globe valves and quick-opening gate valves
that are used to regulate liquid flow in pipes.
Multiblade (louvered) dampers operate by means of the same
principal However, instead using a single blade or plate to
control the gas flow, multiblade dampers employ slats that open
and c?ose like Venetian blinds.15 Louvered dampers typically are
used in 5ery large ducts where a one-piece damper blade would be
too difficult to move.
Manually-controlled dampers simply have a handle attached to
ilct -L_y , ,• V, ' - ' *-~ -,^1-^ot- i-Vi^ rr;=!K flow bv hand. If
rod which i
cco c npn
is uSJd Se actuator receives a pneumatic (pressurized air) or
1 - s -
SET*
depe on the combustibles concentration (i.e., percentage of
Ser explosive limit— ^LEL) in the inlet waste gas stream. If
tMs concentration deviates from a predetermined amount ("set
point"? a signal is sent from the measuring device via the
Controller to the automatic damper to increase/decrease the
dilution air flow rate so as to maintain the desired «LEL.
Expansion joints are installed, especially in longer metal
duct runs to allow the ductwork to expand or contract in
sS to thermal stresses. These fittings are of several
s One type, the bellows expansion joint, consists of a
of flSxible metal (e.g., 304 stainless steel) that is
to each of two duct ends, connecting them. As the
empeatSrfof the duct increases, the bellows compresses; as the
duct temperature decreases, the bellows expands.
the
t
coated ??be?glass cloth is needed to accommodate temperatures of
, 000°F.17
The last component to consider is the ductwork support
system However, it is far from being the least important. As
10-12
-------
the SMACNA (Sheet Metal and Air Conditioning Contractors
National Association) HVAC Duct Construction Standards manual
states "The selection of a hanging system should not be taken
lightly, since it involves not only a significant portion of the
erection labor, but also because [the erection of] an inadequate
hanging system can be disastrous." As a rule a support should be
provided for every 8 to 10 feet of duct run.18 Ductwork can be
suspended from a ceiling or other overhead structure via hangers
or supported from below by girders, pillars, or other supports.
A suspension arrangement typically consists of an upper
attachment, a hanger, and a lower attachment. The upper
attachment ties the hanger to the ceiling, etc This can be a
concrete insert, an eye bolt, or a fastener such as a rivet or
nailed pin. The hanger is generally a strap of galvanized steel,
round steel rod, or wire that is anchored to the ceiling by the
upper attachment. The type of hanger used will be_dictated by
the duct diameter, which is proportional to its weight per lineal
foot For instance, wire hangers are only recommended for duct
diameters up to 10 inches. For larger diameters (up to 36
inches), straps or rods should be used. Typically, a strap
hanqer is run from the upper attachment, wrapped around the duct,
and secured by a fastener (the lower attachment). A rod hanger
also extends down from the ceiling. Unlike strap hangers, they
are fastened to the duct via a band or bands that are wrapped
around the circumference. Duct of diameters greater than 3 feet
should be supported with two hangers, one on either side of the
duct and be fastened to two circumferential bands, one atop and
one below the duct.19 Moreover, supports for larger ductwork
should also allow for both axial and longitudinal expansion and
contraction, to accomodate thermal stresses.
10.2.3 Stacks
Short stacks are installed after control devices to disperse
the exhaust gases above ground level and surrounding buildings.
As opposed to "tall" stacks, which can be up to 1000 feet high,
short stacks typically are no taller than 120 feet.
Certain packaged control devices come equipped with short
("stub") stacks, with heights ranging from 30 to 50 feet. But if
such a stack is neither provided nor adequate, the facility must
erect a separate stack to serve one or more devices.
Essentially, this stack is a vertical duct erected on a
foundation and supported in some manner. For structural
stability, the diameter of the stack bottom is slightly larger ^
than the top diameter, which typically ranges from 1 to 7 feet.'
A short stack may be fabricated of steel, brick, or plastic
(e g fiberglass-reinforced plastic, or FRP). A stack may be
lined'or unlined. The material selection depends on the physical
10-13
-------
The
and chemical properties of the gas stream, such as corrosiveness
and acidity, as well as the temperature differential between the
qas stream and the ambient air. Liners of stainless steel,
brick or FRP usually are used to protect the stack against
damage from the gas stream. They are much easier and less
expensive to replace than the entire stack. Alternatively, the
interior of an unlined stack may be coated with zinc _
(galvanized), aluminum, or another corrosion-resistant material,
but a coating does not provide the same protection as a liner and
does not last as long.22
Short stacks are either self-supporting (free-standing),
supported by guy wires, or fastened to adjacent structures. T
type of support used depends on the stack diameter, height and
weigh?, the wind load, local seismic zone characteristics, and
other factors.
Auxiliary equipment for a typical stack includes an access
door a sampling platform, ladders, lightning protection system
and aircra^ warning lights. The access door allows for removal
of any accumulated materials at the bottom of the stack and
provides access to the liner for repair or replacement. Local
and state air pollution control regulations also may require the
permanent installation of sampling platforms for use during
periodic compliance tests, while ladders are used both _ during
stack sampling and maintenance procedures. The lightning
protection system is needed to prevent damage to the stack and
immediate surroundings during electrical storms^ Lastly,
aircraft warning lights are required by local aviation
authorities.23 Altogether, these auxiliaries can add a large
amount to the base stack cost.
10.3 Design Procedures
As stated above, a hood, ductwork, and a stack are key
elements in any air pollution control system. Because each of
elements is different, both in appearance and function
hs
tee eemen ,
each must be designed separately. But at the same time, these
Cements comprise^ system, which is governed by certain physical
law^ that serve to unite these elements in
"coLon cause". Thus, before the individual design procedures
for hoods, ductwork, and stacks are described, ventilation
fundamentals will be presented. These fundamentals will cover
basic fluid flow concepts and how they may be applied to air
pollution control ventilation systems. Nonetheless these
concepts will be given as straightforwardly as possible, with the
aim of making the design parameters easy to understand and
compute .
10-14
-------
10.3.1 Design Fundamentals
10.3.1.1 The Bernoulli Equation
The flow of fluids in any hood, duct, pipe, stack, or other
enclosure is governed by a single relationship, the familiar
Bernoulli equation. Put simply and ideally, the Bernoulli
equation states that the total mechanical energy of an element of
flowing fluid is constant throughout the system. This includes
its potential energy, kinetic energy, and pressure energy.
However as no system is ideal, the Bernoulli equation must be
adjusted to take into account losses to the surroundings due to
friction. Gains due to the energy added by fans, pumps, etc.,
also must be accounted for. For a pound mass (IbJ of fluid
flowing in a steady-state system the adjusted Bernoulli equation
is:24
Jvdp + Az(g/gc) + A(u2)/2gc = W - F (10.1)
where: v = specific volume of fluid (ft /lbm)
p = static pressure—gauge (lbf/ft )
z = height of fluid above some reference point (ft)
u = fluid velocity through duct, hood, etc. (ft/sec)
g = gravitational acceleration (ft/sec2)
gc = gravitational constant (32.174 ( [lbm-ft/sec ]/lbf)
W = work added by fan, etc. (ft-lbf/lbm)
F = energy lost due to friction (ft-lbf/lbm)
Each of the terms on the left .hand side of equation 10.1
represents an energy change to a pound mass of fluid between two
locations in the system—points "1" and "2". The work (W) and
friction (F) terms denote the amounts of energy added/lost
between points 1 and 2.
Note that the units of each term in equation 10.1 are "ft-
lbf/lbm, " energy per unit mass. In the English system of units,
"lbf" and "lbm" are, for all intents, numerically equivalent,
since the ratio of the gravitational acceleration term (g) to the
gravitational constant (gc) is very close to 1. In effect,
therefore, the equation units are "feet of fluid" or "fluid head
in feet". In air pollution control situations, the fluid often
has the properties of air. That is because the contaminants in
the waste gas stream are present in such small amounts ^ that the
stream physical properties approximate those of pure air.
Because air is a "compressible" fluid, its specific volume
is much more sensitive to changes in pressure and temperature
than the specific volume of such "incompressible" fluids as
water. Hence, the "vdp" term in the equation has to be-
integrated between points 1 and 2. However, in most air
pollution control ventilation systems neither the pressure nor
10-15
-------
the temperature changes appreciably from the point where the
emissions are captured to the inlet of the control device.
Consequently, the specific volume is, for all practical purposes,
constant throughout the ventilation system, and one does not have
to integrate the vdp term. With this assumption, the first term
in equation 10.1 becomes simply:
jvdp = vfdp = vAp (10.2)
Illustration: VOC emitted by an open tank is captured,by a hood
and conveyed, via a blower, through 150 feet of 12-inch diameter
ductwork to a refrigerated condenser outdoors. The blower, which
moves the gas through the hood, ductwork, and condenser, is
located immediately before the inlet to the condenser. Thus, the
entire ventilation system is under vacuum. The stream^
temperature and absolute pressure are 100 °F and approximately 1
atmosphere (14.696 lbr/in2) , respectively. The elevation of the
refrigerated condenser inlet is 30 feet below that, of the tank.
The air velocity at the source is essentially zero, while the
duct transport velocity is 2,000 ft/min. The static gauge
pressure increases from -0.50 in. w.c. (water column) at the
source to 4.5 in. w.c. at the blower outlet. Finally, the
calculated friction loss through the ductwork and hood totals
1 25 in we Calculate the amount of mechanical energy that the
biower adds to the gas stream. Assume that the gas temperature
remains constant throughout.
Solution:
•a* First, develop a factor to convert "inches of water" to "feet
of air":
Feet of air = (Inches of water) (1 ft/12 in) (v,100/vwloo) (10.3)
where- v 100 = specific volume of water @ 100°F = 0.01613 ft3/lbm
v^cT = specific volume of air @ 100°F, 1 atmosphere
Because the system absolute pressure is close to
atmospheric, the waste gas behaves as an ideal gas. Thus, the
specific volume can be calculated from the ideal gas law:
va = RT/pM <10-4)
where: R = ideal gas constant = 1,545 f t- lbf/ (lbm-mole) (°R)
T = absolute temperature of gas = 100 + 460 = 560°R
M = molecular weight of gas (air) =
.28.85 lbm/lbm-mole ^
p = absolute pressure = 2,116 lbf/ft
Substituting, we obtain:
10-16
-------
va = 14.17 ft
Finally, substitution of these values for va and vw into equation
10.3 yields:
Feet of air (@ 100°F, 1 atm.) = 73.207 x Inches of water
cr Compute the changes in the mechanical energy terms and the .
friction losses between the hood inlet (point 1) and the blower
outlet/condenser inlet (point 2):
Pressure: vAp = (4.5 - [-0.50] in. w.c.) (73.207 ft air/in, w.c.)
366.0 ft air
Potential: Az = -30 ft air (point 2 is below point 1)
Kinetic: Au2/2gc = ([2,000 ft/min] / [60 ft/min/1 ft/sec] )2 x
(1/2) (32.174 [lbm-ft/sec2] /lbf)''
= 17.3 ft air
Friction losses: F = 1.25 in. w.c. x 73.207
= 91.5 ft air
isr Substitute above results into equation 10.1 and solve for W,
the fan energy added:
366.0 + (-30) + 17.1 = W - 91.5, or
W = 444.6 ft-lbr/lbm air = 6.07 in. w.c.
To convert the fan energy input, W, to horsepower (hpf) , we
would have to multiply it by the air mass flow rate (lbm/sec),
and divide the result by the horsepower conversion factor, 550
ft-lbf/sec-hp. However, the mass flow rate is just the volume
flow rate (Q, ftVsec) divided by the specific volume:
hpf = W(Q/v.) (1/550) = 0.001818WQ/va (10.5)
(The reader may wish to compare this equation to the fan
horsepower equation in Chapter 3 [page 3-55] of this manual.)
In turn, Q is a function of the duct velocity (ut/ ft/sec) and
duct diameter (Dd, ft) :
Q = u((7TDd2/4) (10.6)
Equation 10.6 applies/ of course, only to circular ducts.
If we combine equations 10.5 and 10.6 and substitute the
inputs for this illustration, we obtain:
10-17
-------
hpf = (444.6) (2,000/60) (7T/4) (1) 2 (1/14 .17 ) (1/550!
= 1.49 hp
Some observations about this illustration:
*r Recall that the precise units for W and the other terms in
equation 10.1 are "ft-lbf/lbm air," which, for convenience have
been shortened to "ft air". Thus, they measure energy, not
length.
«r Compared to the pressure energy and friction terms, the
potential and kinetic energy terms are small. Had_they been
ignored, the results would not have changed appreciably.
«• The large magnitude of the pressure and friction terms clearly
illustrates the importance of keeping_one's^units straight. As
c^hown in step (1) , one inch of water is equivalent to over 73
feet of air However, as equation 10.3 indicates, the pressure
corresponding to equivalent heights of air and water columns
would be the same.
«• The fan power input depends not just on the total "head" (ft
air) reauired but also on the gas flow rate. Also, note that
?he horsepower computed via equation 10.5 is a theoretical value.
It woSld have to be adjusted to account for the efficiencies of
the fan and fan motor. As mentioned in Chapter 3, the fan
efficiency ranges from 40 to 70 percent, while the motor
efficiency is typically 90 percent. These efficiencies are
usually comJineFinto a single efficiency U._fraction), by which
the theoretical horsepower is divided to obtain the actual
horsepower requirement.
10.3.1.2 Pressure: Static, Velocity, and Total
Although it is more rigorous and consistent to express the
Bernoulli equation terms in terms of feet of air (or, precisely,
Bernoulli equ industrial ventilation engineers prefer to
use the units "inches of water column (in. w.c.) " These units
were chosen because, as the above illustration shows, results
expressed in "feet of air" are often large numbers that are _
cumbersome to use. In addition, the total pressure changes in
ventilation systems are relatively small, compared to those in
IfSuid flow systems. Total pressure changes expressed in inches
of mercury would be small numbers which are Dust as awkward to
work with as large numbers. Hence, "inches of water" is a
compromise, as values expressed in this measurement unit
topically range from only 1 to 10. Moreover, practical
measurement of pressure changes is done with water-filled
10-18
-------
manometers.
in the previous paragraph, a new quantity was mentioned
total pressure (TP). Also known as the "impact pre^ure«, the
total pressure is the sum of the static gauge JSP) and velocity
pressures (VP) at any ^oint within a duct, hood, etc., all
expressed in in. w.c.25 That is:
TP = SP + VP (10.7)
where: SP = (cf)vp
VP = (cf)uY2gc
The "cf" in the expressions for SP and TP is the factor for
converting the energy terms from "ft air" to "in. w.c *oth f
standard temperature and absolute pressure (70°F l atmosphere^
(Again, keep in mind that, regardless of what units SP or VP are
expressed in, the actual units are "energy per unit mass".) This
conversion factor would be obtained via rearranging equation
10.3:
cf = in. w.c./ft. air = 12 (vw70/va70)
where: vw70 = specific volume of water at 70°F = 0
Va7o = specific volume of air at 70°F = 1.
(10.8)
Thus:
cf = 0.01436 in. w.c. /ft air
Clearly, "cf" varies as a function of temperature and
pressure For instance, at 100°F and 1 atmosphere, cf = 1/73.207
= 0 01366. Nevertheless, unless noted otherwise, all quantities
henceforth in this chapter will reflect conditions at 70°F and 1
atmosphere.
Conspicuously absent from equation 10.7 is the potential
energy term, "z(g/gc)"- This omission was not inadvertent In
ventilation systems, the potential energy (P.E. ) . is ^^^ sma11
compared to the other terms. (For example, see illustration
above ) The P.E. is, of course, a function of the vertical
distance of the measurement point in question from some datum
level usually the ground. At most, that distance would amount
to no more than 20 or 30 feet, corresponding to a P.E. of
approximately 0.3 to 0.4 in. w.c. Consequently, we can usually
ignore the P.E. contribution in ventilation systems without
introducing significant error.
The static gauge pressure in a duct is equal in all
directions, while the velocity pressure, a function of _ the _
velocity, varies across the duct- face. The duct velocity is
highest at the center and lowest at the duct walls. However, for
air flowing in a long, straight duct, the average velocity (u,)
10-19
-------
approximates the center line velocity (ucl) .2 This is an
important point, for the average velocity is often measured by a
pitot tube situated at the center of the duct.
By substituting for "cf" in equation 10.7, we can obtain a
simple equation that relates velocity to velocity pressure at
standard conditions:
VP = 0.01436u,2/2gc (10.9)
Solving:
u( (ft/sec) = 66.94(VP)"2 (10-1Q)
Or:
u, (ft/min) = 4,016(VP)1/2 (10.11)
Incidentally, these equations apply to any duct, regardless
of its shape.
As Burton describes it, static gauge pressure can be thought
of as the "stored" energy in a ventilation system.. This stored
energy is converted to the kinetic energy of velocity and the
losses of friction (which are mainly heat, vibration, and noise).
Friction losses fall into several categories:
B3" Losses through straight duct
•s- Losses through duct fittings—elbows, tees, reducers, etc.
•5- Losses in branch and control device entries
•s- Losses in hoods due to turbulence, shock, vena contracta
i®" Losses in fans
«3" Losses in stacks
These losses will be discussed in later sections of this
chapter Generally speaking, much more of the static gauge _
pressure energy is lost to friction than is converted_to velocity
pressure energy. It is customary to express these friction
losses (ASPf) in terms of the velocity pressure:
F = ASPf = kVP _ <10-12)
where: k = experimentally-determined loss factor (unitless)
10-20
-------
Alternatively, equations 10.11 and 10.12 may be combined to
express F (in. w.c.) in terms of the average duct velocity, ut
(ft/min):
F = (6.200 x 10-8)kUl2 (10.13)
10.3.1.3 Temperature and Pressure Adjustments
Equations 10.8 to 10.13 were developed assuming that the
waste gas stream was at standard temperature and pressure These
conditions were defined as 70°F and 1 atmosphere (14.696 lbf/in ) ,
respectively While 1 atmosphere is almost always taken as the
standard pressure, several different standard temperatures are
used in scientific and engineering calculations: 32°F, 68 F, and
77°F as well as 70°F. The standard temperature selected varies
according to the industry or engineering discipline in question.
For instance, industrial hygienists and air conditioning
engineers prefer 70°F as a standard temperature, while combustion
engineers prefer 77°F, the standard temperature used in Chapter j
("Thermal and Catalytic Incinerators").
Before these equations can be used with waste gas streams
not at 70°F and 1 atmosphere, their variables must be adjusted.
As noted above, waste gas streams in air pollution control
applications obey the ideal gas law. From this law the following
adjustment equation can be derived:
Q2 = Q,(T2/T,) (P,/P2) (10.14)
where: Q2,Q, = gas flow rates at conditions 2 and 1,
respectively (actual ft3/min)
T2,T, = absolute temperatures at conditions 2 and 1,
respectively (°R)
P2,P, = absolute pressures at conditions 2 and 1,
respectively (atm)
However, according to equation 10.6:
Q = u,(7rDd2/4)
If equations 10.6 and 10.14 were combined, we would obtain:
UQ = utl(T2/T.) (P,/P2) (DJ22/Ddl2) (10.15)
This last expression can be used to adjust u, in any
equation, as long as the gas flow is in circular ducts.
10-21
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10.3.2 Hood Design Procedure
10.3.2.1 Hood Design Factors
When designing a hood, several factors must be considered:2&
B5* Hood shape
BS* Volumetric flow rate
"3" Capture velocity
cs* Friction
Each of these factors and their interrelationships will be
explained in this section.
As discussed in section 10.2.1, the hood shape is determined
by the nature of the source being controlled. This includes such
factors as the temperature and composition of the emissions, as
well as the dimensions and configuration of the emission stream.
Also important are such environmental factors as the velocity and
temperature of air currents in the vicinity.
The hood shape partly determines the volumetric flow rate
needed to capture the emissions. Because a hood is under
negative pressure, air is drawn to it from all directions.
Consider the simplest type of hood, a plain open-ended duct.
Now, envision an imaginary sphere surrounding the duct opening.
The'center of this sphere would be at the center of the duct
opening, while the sphere radius would be the distance from the.
end of the duct to the point where emissions are captured. The
air would be drawn through this imaginary sphere and into the
duct hood. Now, the volume of air drawn through the sphere would
be the product of the sphere surface area and the hood capture
velocity, uc:29
Q = uc(47rx2) (10.16)
where: x = radius of imaginary sphere (ft)
Equation 10.16 applies to a duct whose diameter is small
relative to the sphere radius. However, if the duct diameter is
larger, the capture area will have to be reduced by the cross-
sectional area of the duct (Dd) , or:
Q = uc(47rx2 - 7TD//4) (10.17)
Similarly, if a flange were installed around the outside of
the~~duct end, the surface area through which the air was
drawn and the volume flow rate—would be cut in half. That
occurs because the flange would, in effect, block the flow of air
from points behind it. Hence:
10-22
-------
Q =
From these examples, it should be clear that the _ hood shape
has a direct bearing on the gas flow rate drawn into it. But
equations 10.16 to 10.18 apply only to hoods with spherical flow
patterns. To other hoods, other flow patterns
apply— cylindrical, planal, etc. We can generalize this
relationship between volumetric flow rate and hood design
parameters as follows:
Q-fCu,. x. Sh) . „ <10-191
where: "f (...)" denotes "function of...
"Sh" indicates hood shape factors
u, = design velocity — capture, face, slot
Table 10 1 lists design equations for several commonly used
hood shapes. 'AS this table shows, Q is a function of x. the .hood
shape, and, in general, the capture velocity (uc) But in one
case (booth hood), the design velocity is the hood face velocity
(uf) And in the case of slotted side-draft and back-draft
hoods, the slot velocity (us) is the design velocity. In
reality, both the hood face and slot velocities are the same as
each measures the speed at which the gas passes through the hood
inlet opening (s) .
When gas enters a hood, there is mechanical energy loss due
to friction This friction loss is calculated using equations
10 1 and 10.2, assuming that the potential energy contribution
from gravity, Az (g/g.) . and the work added to the system, W, are
both zero. Thus:
vp2 - vp, + u22/2gc - u,2/2gc = -
(10.20)
Replacing these terms with the corresponding ones from
equations 10.7 and 10.12, we obtain:
SP2 - SP, + VP2 - VP, = - Hc = - khVP2 (10.21)
where: SP, = static gauge pressure at point i (in. w.c.)
VP] = velocity pressure at point i (in. w.c.)
H, = hood entry loss (in. w.c.)
k|] = hood loss factor (unitless)
In this equation, subscript 1 refers to a point just outside
the hood face. Subscript 2 denotes the point in the duct, just
downstream of the hood, where the duct static pressure, SP2 or
SP and the duct transport velocity, u2 or ut, are measured. At
point 1 the hood velocity pressure, VP,, is essentially zero, as
the air velocity there is negligible. Moreover, the static gauge
10-23
-------
Table 10.1 Design Equations, Loss Factors, and Coefficients of
Entry for Selected Hood Types*
Hood Type
Duct end
(round)
Flanged duct
end (round)
Free-standing
slot hood
Slot hood
w/sides, back
Tapered hood
Booth hood with
tapered take-off
duct (round)
Canopy hood
Canopy hood
w/insert
Dip tank hood
(slotted)
Paint booth
hood
Design
Equation*
Q = 4?rx2uc
Q = 27TX2UC
Q = 27rxLuc
Q = 0 . 5 7rxLuc
Q = 27TXUC
Q = uA,
Q = 1.4Pxuc
Q = 1.4Pxuc
Q = 125A,
Q = lOOA,,
Loss Factor
-------
pressure, SP,, will be zero, as the absolute pressure at point 1
is assumed to be at one atmosphere, the reference pressure.
After these simplifications are made, equation 10.21 can be
rearranged to solve for the hood loss factor (kh) :
kh = (-SPh/VP2) - 1 (10.22)
At first glance, it appears that kh could be negative, since
VP is always positive. However, as the air entering the hood is
under a vacuum created by a fan downstream, SPh must be negative.
Thus, the term "-SPh/VP2" must be positive. Finally, because the
absolute value of SPh is larger than VP2, kh > 0.
The hood loss factor varies according to the hood shape. It
can range from 0.04 for bell mouth hoods to 1.78 for various_
slotted hoods. A parameter related to the hood loss factor is
the coefficient of entry (ce).30 This is defined as:
ce = {i/d+kj}
1/2 (10.23)
c depends solely on the shape of the hood, and may be used to
compute kh and related parameters. Values of kh and ce are listed
in Table 10.1.
Illustration: The static gauge pressure, SPh, is -1.75 in. w.c.
The duct transport velocity (u,) is 3,500 ft/min. Calculate the
loss factor and coefficient of entry for the hood. Assume
standard temperature and pressure.
Solution: First, calculate the duct velocity pressure. By
rearranging equation 10.11 and substituting for ut, we obtain:
VP = (ut/4,016)2 = (3,500/4,016)2 = 0.76 in. w.c.
Next, substitute for VP in equation 10.22 and solve:
kh = (_SPh/VP) - 1 = (- [-1-75J/0.76) - 1 = 1.30.
Finally, use this value and equation 10.23 to calculate the
coefficient of entry:
ce = (l/d + 1.30) }1/2 = 0.66.
Hood design velocities are listed in Table 10.2. Three
kinds of velocities are shown: (1) capture (defined in Section
10 2.1), (2) face, and (3) slot. As stated in Section 10.2.1,
the capture velocity is the air velocity induced by the hood to
capture contaminants emitted at some distance from the hood
inlet. The face velocity is the average velocity of the air
10-25
-------
nassina through the hood inlet (face) . A similar parameter is
?he sSot vSocity, which is the average air velocity through the
hSod s?ot openings, whose area is only a fraction^ the entire
hood face area. Consequently, the slot velocity is usually much
higher than the face velocity.
Note that these velocities range from 50 to 100 ft/min (tank
and decreasing hoods) to 2,000 ft/min, the recommended slot
ve?ocf?y for slotted side-draft/back-draft hoods. As a reference
noint the velocity of air in industrial operations due to
Serial mixing alone is 50 ft/min. Thus, hood design velocities
musTJxceed this value if effective capture is to occur.
Two other velocities are also discussed in the industrial
the vlenum velocity and the transport velocity. The first
velo?i5TSf the gas stream as it passes through the tapered
a
i-o.
rrucTal paramSte? in determining the duct diameter, the static
pressured™ and the sizes of the system fan and fan motor.
uSr more on transport velocity, see Section 10.3.3.)
10.3.2.2 Hood Sizing Procedure
As with manv control devices and auxiliaries, there are
i^
hood cost.
This method does yield reasonably accurate hood co
rather it did. Unfortunately, the labor cost data are
—- vintage—^hich makes them unescalatable. (The
10-26
-------
use, especially if calculations are made by hand.
A simpler sizing method— yet one sufficiently accurate for
smsrs.1- Ls-ss-^r ~^iSa '
hood inlet area, can be correlated against the fabricated
is needed:
e^ Hood type
«• Distance of the hood face from source (x)
•ar Capture (uc) , face (uf) , or slot velocity (u.)
us- Source dimensions (for some hood types) .
As the equations in Table 10.1 indicate, these same ^
narameLrs are ?he ones that are used to determine the volumetric
f?ow rate (^through the hood and ductwork. With most control
devices and auxiliaries being sized, Q is given. For hoods,
however, Q usually must be calculated.
Illustration: A circular canopy hood^is being used £o capture
emissions from a chromium electroplating tank. The hood face is
6 feet above the tank, an 8- foot diameter circular vessel. The
capture velocity for this example is 200 ft/min. _ Assuming that
the tank and surroundings are at standard conditions calculate
the Squired volumetric flow rate drawn into the hood, the hood
face area, and the hood face velocity.
Solution: Obtain the canopy hood equation from Table 10.1:
Q = 1.4Pxuc (10"24)
where: P = perimeter of tank (ft)
x = distance of hood above tank (ft)
Uc = capture velocity (ft/min)
Because the tank is circular, P = 7r(8) = 25.1 ft.
Therefore:
Q = (1.4) (25.1) (6) (200) = 42,200 ft3/min.
For this type of canopy hood, the hood diameter is 40%
greater than the tank diameter (hence, the "1.4" factor in
equation 10.24) . Thus:
A, = (TT/4) ([1-4] [8] )2 = 98.5 ft2
Finally, the hood face velocity (uf) would be:
10-27
-------
Table 10.2 Hood Design Velocities*
1-
Operat ion/Hood Type
Tanks, degr easing
Drying oven
Spray booth
Canopy hood
Grinding, abrasive
blasting
Slot hood
^^^=^^=======
Velocity Type
Capture
Face
Capture
Capture
Capture
Slot
Velocity Range
(ft /min)
50 - 100
75 - 125
100 - 200
200 - 500
500 - 2,000
2, 000
** Reference: Burton, D. Jeff. Industrial Ventilation Work
Book. Salt Lake City: DJBA, Inc. 1989.
10-2!
-------
U =
= 42,200/98.5 = 428 ft/min.
in this example, note that the hood face velocity "higher
than the capture velocity. This is logical, gxven the fact that
the hood inlet area is smaller than the area through which the
tank fumes are being drawn. The face velocity for some hoods xs
even higher. For example, for slotted hoods it xs at least 1 000
ft/min % In fact, one vendor sizes the openings xn his slotted
hoods so as to achieve a slot velocity equal to the duct
transport velocity.35
10.3.3 Ductwork Design Procedure
The design of ductwork can be an extremely complex
undertaking. Determining the number, placement, and dimensxons
of ductwork components—straight duct, elbows, tees, dampers,
etc can be tedious and time-consuming. However, for purposes
of making study-level control system cost estimates, such
involved design procedures are not necessary. Instead, a much
simpler ductwork sizing method can be devised.
10.3.3.1 Two Ductwork Design Approaches
There are two commonly used methods for sizing and pricing
ductwork. in the first, the total weight of duct is computed
from the number and dimensions of the several components. Next,
this weight is multiplied by a single price (in $/lb) to obtain
the ductwork equipment cost. To determine the ductwork weight,--
one needs to know the diameter, length, and wall thickness of
every component in the system. As stated above, obtaining these
data can be a significant effort.
The second method is a variation of the first. In this
technique, the ductwork components are sized and priced
individually. The straight duct is typically prxced as a
function of length, diameter, and wall thickness, as well as, of
course the material of construction. The elbows, tees, and
other fittings are priced according to all of these factors,
except for length. Other variables, such as the amount and type
of insulation, also affect the price. Because it provides^more
detail and precision, the second method will be used in this
chapter.
10.3.3.2 Ductwork Design Parameters
Again the primary ductwork sizing variable are length,
diameter, and wall thickness. Another parameter is the amount of
insulation required, if any.
10-29
-------
•5" Length: The length of ductwork needed with an air pollution
control system depends on such factors as the distance of the
source from the control device and the number of directional
changes required. Without having specific knowledge of the
source layout, it is impossible to determine this length
accurately. It could range from 20 to 2,000 feet or more. It is
best to give the straight duct cost on a $/ft basis and let the
reader provide the length. This length must be part of the
specifications of the emission source at which the ductwork is
installed.
«• Diameter: As discussed in Section 10.2.2., circular duct is
preferred over rectangular, oval, or other duct shapes.
Therefore:
A, = 7rDd2/4 (10.25)
where: A,, = cross-sectional area of duct (ft2)
Dd = duct diameter (ft)
The duct cross-sectional area is the quotient of the volumetric
flow rate (Q) and the duct transport velocity (u,) :
A, = Q/u, (10.26)
Combining equations 10.25 and 10.26 and solving for Dd:
Dd = 1.128(Q/ut)1/2 (10.27)
As Q is usually known, the key variable in equation 10.27 is
the duct transport velocity. This variable must be chosen
carefully. If the u, selected is too low, the duct will be
oversized and, more importantly, the velocity will not be high
enough to convey the particulate -matter in the waste gas stream
to the control device. However, if u, is too high, the static
pressure drop (which is proportional to the square of ut) will be
excessive, as will be the corresponding fan power consumption.
Cost is also a consideration when determining the optimum
duct diameter. The equipment cost increases with increasing duct
diameter. However, the fan power cost changes inversely with
diameter. Nonetheless, for study-estimating purposes, the
optimum duct diameter does not have to be determined. It is
sufficient to calculate the duct diameter merely_by using the
transport velocity values contained in this section.
The transport velocity typically varies from 2,000 to 6,000
ft/min, depending on the waste gas composition. The_ lower duct
velocity would be adequate for a waste gas containing gaseous
pollutants or very fine, light dusts, while the higher velocity
would be needed to convey a stream with a large quantity of
10-30
-------
metals or other heavy or moist materials. The following
velocities may be used as general guidance:
Material(s) Conveyed
Minimum Transport Velocity
(ut, ft/min)
Gases- very fine, light dusts
2,000
Fine dry dusts and powders
3,000
Average industrial dusts
3,500
Coarse dusts
4,000 - 4,500
Heavy or moist dust loading
Table 10.3 supplements these values with recommended duct
velocities for a variety of conveyed materials.
*r wall thickness: The wall thickness of a duct depends on
several factors—internal pressure, diameter, material of
fabrication, and other structural parameters. Nonetheless, duct
of a given diameter can be fabricated of a range of wall
thicknesses, and vice-versa. For instance, 24-in. diameter 304
stainless steel "fully-welded longitudinal seam duct" is
fabricated in thicknesses ranging from 22 to 14 gauge (0.0313 to
0.0781 in.). This same range of gauges is used with duct
diameters ranging from 3 to 36 in.
Note that the gauge number decreases with increasing wall
thickness This measure, which is traditionally used in the
metal fabricating industries, is more convenient to deal with
than the thickness expressed in inches, as the latter are usually
small numbers less than 0.25. Moreover, the gauge number varies
according to the metal used—carbon steel (galvanized or
nongalvanized), stainless steel, or aluminum. Gauges for these
metals are given in Table 10.4 for a wide range of nominal
thicknesses.
The gauge measure is not used with plastic duct, as the wall
thickness is typically expressed in inches. In any event, the
wall thickness usually does not need to be known to estimate duct
cost, as this parameter is already accounted for in the cost
equations. (See Section 10.4.)
«r Insulation: As discussed in Section 10.2.2., insulation can be
either installed on the outer surface of ductwork or the ductwork
itself can be fabricated -with built-in insulation. In the-first
case the amount of insulation required will depend on several
heat'transfer variables, such as: the temperature, velocity,
10-31
-------
composition, and other properties of the waste gas; the cimbient
temperature; the duct diameter, wall thickness, arid thermal
conductivity; and the desired surface ("skin") temperature.
Determining these variables involves making a series of complex
calculations that, while well-established, are beyond the scope
of this chapter. Such standard references as Perry's Chemical
Engineers' Handbook and Plant Design and Economics for Chemical
Engineers present these calculations, as do heat transfer
^i? ^Q
bibl iographies . J*'JV
The second approach is to select pre- insulated ductwork. As
mentioned previously, it can be equipped with any type arid _
thickness of insulation. However, 1, 2, or 3 inches is typical.
(Prices for these are presented in Section 10.4.)
10.3.3.3 Ductwork Pressure Drop
As mentioned in Section 10.3.1, ventilation system energy
losses due to friction are traditionally computed as fractions of
the velocity pressure, VP. In most cases, equation 10 1,. can be
used to estimate these losses. Technically, though, these
equations apply only to those regions in the ventilation system
where there are no changes in the velocity pressure (i.e where
the duct diameter is constant) . These regions would include
straight duct, hoods, and such fittings as couplings and simple
elboii But, with tees, wyes, and other divided flow fittings
the velocity-^md velocity pressure^re not constant between the
fitting inlet and outlet. The corresponding friction loss (Fb)
is a function of both the upstream (inlet) and branch VP's, as
the following equation indicates:
Fb = VPu(kb-l)
VP (10.28)
b
where: VPU, VPb = upstream and branch velocity pressures,
respectively (in. w.c.)
kh = branch loss coefficient
However divided flow fittings generally are not used with simple
pollution control ventilation systems, except in those cases
where a tee might be needed, say, for purposes of adding dilution
6
air.
As any fluid mechanics textbook would attest, the friction
loss for ductwork is a complex function of several variables:
« Divided flow fittings are needed with more-complex _
control systems that collect waste gases from several emission
points. The design of such ventilation systems is beyond the
scope of this chapter, however.
10-32
-------
Table 10.3 Minimum Duct Velocities for Selected Materials-
Material
Aluminum dust (coarse)
Brass turnings
Cast iron boring dust
Clay dust
Coal dust (powdered)
Cocoa dust
Cotton dust
Flour dust
Foundry dust
Grain dust
Lead dust
Limestone dust
Magnesium dust (coarse)
Metal turnings
Plastics dust (buffing)
Rubber dust
Silica dust
Soap dust
Soapstone dust
Spray paint
Starch dust
Stone dust
Tobacco dust
Minimum Transport Velocity
(ft/min)
4,000
4,000
4,000
3,500
4,000
3,000
3,000
2,500
3,000 - 5,000f
2,500 - 3,000
4,000
3,500
4,000
4,000 - 5,000
3,000
2,500 (fine) - 4,000 (coarse)
3,500 - 4,500
3,000
3,000
2,000
3, 000
3,500
3,500
§ Reference: Burton, D. Jeff. Industrial Ventilation
Book. Salt Lake City: DJBA, Inc. 1989.
1 Transport velocity varies with foundry operation.
10-33
-------
Table 10.4 Wall Thicknesses of Steel and Aluminum Duct5
Gauge
Number
28
26
24
22
20
18
16
14
12
10
Nominal Thickness (inches)
Carbon Steel
Galv*
0.0187
0.0217
0.0276
0 .0336
0.0396
0.0516
0.0635
0 . 0785
0.1084
0.1382
Nongalv*
0.0149
0.0179
0.0239
0.0299
0.0359
0.0478
0,0598
0.0747
0.1046
0.1345
Stainless Steel
(304 or 316)
0.0156
0.0188
0.0250
0.0313
0.0375
0.0500
0.0625
0.0781
0.1094
0.1406
Aluminum
3003-H14f
0.025
0.032
0.040
0.050
0.063
0.080
0.090
§ Reference- Engineering Design Reference Manual for Supply
Air Handling Systems. Groveport, OH: United McGill Corporation.
1992.
f To provide equivalent strength and stiffness, the nominal
thickness of aluminum is approximately 150% of the nominal
thickness of galvanized carbon steel of the same gauge.
* Galvanized and paintable galvanized carbon steel.
' Nongalvanized carbon steel.
10-34
-------
duct diameter and length, transport velocity, and gas viscosity
and density. Specifically, the Darcy-Weisbach and Colebrook
equations are typically used to make this calculation, the latter
being used to compute the .Reynolds number.41 Traditionally, the
friction loss has been obtained from a nomograph or, more _
recently, computer programs. A typical nomograph is found in
Burton & Also, to simplify the calculation, empirical equations
have been derived for certain kinds of commerically-available
ductwork. For instance, to estimate the friction loss per 100 ft
(F/100 ft) at standard conditions for round, spiral, galvanized
ductwork having 10 joints per 100 ft, use the following
equation:43
Fd/100 ft = 0.136(l/D)1-I8(u,/l,000)1-8 (10.29)
where: Dd = duct diameter (ft), and: 0.25 <. Dd < 5
Clearly this equation provides the total friction loss, not the
loss factor (k) However, the reader may compute k for a given
diameter (Dd) and flow rate (Q) by simply dividing the equation
10.29 results by VP and multiplying by 100.
To estimate the friction loss for other duct materials, _
multiply the value from equation 10.29 b£ a roughness correction
factor, approximate values of which are:
Material
Roughness Correction Factor_
Non-spiral-wound galvanized
0.9
Fiberglass (smooth finish)
0.8
ABS and PVC plastic
0.8
Concrete
1.4
Corrugated flex duct
2.3
Loss factors for fittings have also been compiled, based on
experimental data. Mainly of interest are those for 90° elbows
arguably the most commonly used fitting in air pollution control
systems. The "k90" values for elbows vary according to the
diameter and radius of curvature, which is expressed as a
multiple of the elbow diameter. Typical ranges of these values
are as follows:45
10-35
-------
Radius of
0.
1.
1.
1.
2.
2.
Curvature
50
00
25
50
00
50
Friction
0
0
0
0
Loss Factor (kgo)
0
0
.30
.27
.24
.22
.80
.35
- 0
- 0
- 0
- 0
.55
.39
.27
.24
As these values indicate, the higher the radius of
curvature, the lower the friction loss. This stands to reason,
as the higher the radius of curvature, the more gradually the gas
stream changes direction. For an elbow having of angle less than
90°, multiply the above k<,0 value by an adjustment factor (0/90),
so that:
k, = (0/90)k90
where: k, = loss factor for 6 < 90°
;i0.30)
Illustration: A control device at a cosmetic factory is connected
to a source by 250 feet of round spiral duct. The duct run
includes three 90° elbows and two 45° elbows, each with a 1.50
radius of curvature. The volumetric flow rate (Q) of the waste
gas (which contains entrained face powder) is 15,000 ft3/™in at
standard conditions. Calculate the friction loss for the
ductwork.
Solution: Because the material being conveyed in the ductwork ^
(face powder) is light, an appropriate transport velocity (ut) in
this case is 2,000 ft/min. (See text table above.) Upon
substituting this value and the volumetric flow rate into
equation 10.27 we obtain the duct diameter (Dd) :
Dd = 1.128 (15,000/2, OOO)0'5 = 3.09 ft
Next, substitute the diameter and velocity into equation 10.29 to
compute the straight duct friction (static pressure) loss, Fd:
F, = 0.136(1/3.09)118(2,000/1,000)18(250/100:
= 0.313 in. w.c.
The 250/100 factor in this expression adjusts the friction loss
from 100 feet (the basis of equation 10.29) to 250 feet (the
length of the duct system in this illustration).
10-36
-------
The rest of the friction loss occurs through the five elbows
(three 90°, two 45°) , each with a 1.50 radius of curvature.
These losses (Fe) are computed via equation 10.12:
Fe =
where: VP = (2 , 000/4 , 016)2 (equation 10.11, rearranged)
= 0.248 in. w.c.
For the 90° elbows, ks = k90 = 0.33 (average of table range), and:
Fe = 3 x 0.33(0.248) = 0.246 in. w.c.
For the 45° elbows, ks = (45/90)k90 = 0.165 (equation 10.30),
and:
Fe = 2 x 0.165(0.248) = 0.0818 in. w.c.
The total friction loss is, therefore:
F = 0.313 + 0.246 + 0.0818 = 0.641 in. w.c.
From this illustration, two observations may be made: (1)
the static pressure loss through the straight duct is not large,
even at this length (250 ft.) and (2) the losses through the
elbows—which total 0.328 in. w.c.-— are larger than the straight
duct loss. Though it may be tempting to neglect fittings losses
for the sake of expediency, doing so can cause a significant
underestimation of the ventilation system static pressure loss.
10.3.4 Stack Design Procedures
As with ductwork, the design of stacks involves a number of
stream, structural, and site- specif ic parameters. • These
include :
•a* waste gas variables: inlet volumetric flow rate, temperature,
and composition;
«sr Site-specific data: elevation above sea level, ambient
temperature fluctuations, topographic and seismic data,
meteorological records, and building elevations and layout ;
«S" Structural parameters: thickness of stack wall and liner,
location of breeching opening, type of supports, load _ capacity of
foundation, modulus of resistance, and natural vibration
frequency.
10-37
-------
Fortunately, for study cost-estimating purposes, the only
two stack design parameters that need to be determined are: (1)
the stack diameter and (2) the stack height. The other variables
(e.g., wall thickness) are incorporated into the equipment cost
correlations. The stack diameter is relatively easy to
determine, as it depends primarily on waste stream conditions.
The stack height is more difficult to arrive at, as it is
influenced by several site-specific variables. Nonetheless,
ample guidance has been developed to allow the estimator to
determine an acceptably accurate stack height.
10.3.4.1 Calculating Stack Diameter
Because most stacks have circular cross - sections, the stack
diameter (Ds, ft) can be calculated via the duct diameter formula
(equation 10.27):
Ds = 1.128(Qe/uc)"2 (10.31)
where: uc = stack exit velocity (ft/min)
Qc = exit volumetric flow rate (actual ft /min)
It should be noted that the stack diameter in this formula
is measured at the stack exit, not at the entrance. That is
because, for structural reasons, the diameter at the bottom of
the stack typically is larger than the top diameter. Also note
that the stack exit velocity does not necessarily equal the duct
transport velocity. Finally, Qc may be different from the
volumetric flow rate used to size the ductwork. Because the_
stack always follows the control device, the flow rate entering -
the device may not equal the flow rate entering the stack, either
in standard or actual ft3/min terms. For instance, in a thermal
incinerator, the outlet standard waste gas flow rate is almost
always higher than the inlet flow rate due to the addition of
supplemental fuel.
The stack exit velocity, ue, affects the plume height, the
distance that the plume rises above the top of the stack once it
exits. In a well-designed stack, uc should be 1.5 times the wind
speed. Typically, design exit velocities of 3,000 to 4,000
ft/min are adequate.48 This range corresponds to wind speeds of
34 to 45 mi/hr.
10.3.4.2 Calculating Stack Height
Estimating the stack height is more difficult than
calculating the stack exit diameter. The stack height depends on
several variables: the height of the source; the scack exit
velocity; the stack and ambient temperatures; the height, shape,
and arrangement of the nearby structures and terrain; and the
composition of the stack outlet gas. Some of these variables are
10-38
-------
straightforward to determine, while others (such as the
dimensions and layout of nearby structures) are difficult_to
determine without performing on-site modeling and monitoring
studies.
This height has two components: the height of the stack
itself (H,) and the plume rise height (Hpr) . Together these _
components comprise the effective stack height (He) . That is:
TT _ w , u (10.32)
nc ~ "s pr
However, the cost of the stack is a function of Hs alone.
(See Section 10.4.) As discussed above, the plume rise is a
function of the stack exit velocity. It also depends on the
temperature differential between the stack gas and the ambient
air. Specifically, a 1°F temperature difference corresponds to
approximately a 2.5-ft. increase in H,,,.49
For those sources subject to State Implementation Plans
(SIPs) , the stack height (H,) should be determined according to
"good engineering practice" (GEP). GEP is defined as "the height
necessary to insure that emissions from the stack do not result
in excessive concentrations of any air pollutant in the immediate
vicinity of the source as a result of atmospheric downwash,
eddies, or wakes which may be created by the source itself,
nearby structures, or nearby terrain obstacles."50 In this
respect, GEP establishes the maximum allowable stack height^
credit for purposes of calculating the ambient air quality impact
of the emitting source. A source may build a stack to any
height, but only a certain amount of stack height will be allowed
in determining environmental impacts.51
For stacks constructed after January 12, 1979, the GEP stack
height shall be the greater of: (1) 65 meters (213 ft); (2) the
height demonstrated by an approved fluid model or field study
that ensures that stack emissions do not cause excessive
pollutant concentrations from atmospheric downwash, wakes, eddy
effects., etc; or (3) the height determined by the following
equation:52
Hs = Hb + 1.5L (10.33)
where: Hs = GEP stack height, measured from the ground level
elevation at the stack base (ft)
Hb = height .of nearby structure(s) measured from this
ground level elevation (ft)
L = lesser dimension (height or projected width of
nearby structure(s))
10-39
-------
10.3.4.3 Calculating Stack Draft
As discussed previously, waste gas flowing through hoods and
ductwork loses static pressure due to friction. In the case of
stacks, however, the gas stream can actually gain static
pressure, as a result of stack draft, which is the _ draft created
by the stack gas-ambient air temperature differential. StacK
draft (SPS, in. w.c.) can be calculated as follows:
SPS = 0.034(HS - HJIKI/T^ - 1/TJ (10.34)
where: Hbr = height of stack breeching (inlet duct
connection)
above stack base (ft)
H = barometric pressure (in. w.c.)
Tamh = ambient temperature (°R)
T™' = average stack gas temperature (°R)
Illustration: The waste gas from a thermal incinerator has an
outlet flow rate and temperature of 21,700 actual ft/min._ and
550 OF respectively. The maximum wind speed in the vicinity is
42 mi/hr, while the stack exit and ambient temperatures are 450 F
and 70°F, in turn. The barometric pressure is 1 atm 23.92 in.
Ho) The incinerator is near a 35-ft tall brick building while
the ""projected width" of an adjacent building is 40 ft. F°r.a
stack to disperse the incinerator off gas, calculate the required:
(1) exit velocity, (2) diameter, (3) height, and (4) dratt.
Solution:
** E2cit_veiQcity.: According to the above guideline, the velocity
should be 1.5 times the wind speed, or:
ue = 1.5 x 42 mph x 88 fpm/mph = 5,540 ft/min.
«5> stack diameter: The exit volumetric flow rate is measured at
the stack exit temperature, namely 450°F. However, the above
flow rate was measured at 550°F, the incinerator outlet
temperature. Correcting to the stack exit temperature, ue
obtain:
Qe = 21,700 x (450 + 460)/(550 + 460) = 19,600 actual
f t3.
Substituting this value into equation 10.31:
Ds = 1.128(19, 600/5, 540)1/2 = 2.12 ft.
«sr stack height: As a first approximation, estimate the GEP stack
height from equation 10.33, where the variables Hb and L are 35
ft and 40 ft, respectively:
10-40
-------
H, = 35 + 1.5 (40) = 95 ft.
Clearly, this Hs is less than the GEP maximum height (213 ft) , so
it will be used in this example.
ts- Stack draft: All of the inputs needed to compute the stack
draft via equation 10.34 are known except the stack breeching
height Hh However, a minimum of 5 ft is recommended for this
parameter."54 This value will be used in this calculation. Also,
the average stack temperature is:
TM = (450 + 550)/2 + 460 = 960°R.
Finally, the barometric pressure expressed in inches of water is:
II = 29.92 in. Hg x 13.6 in. water/in. Hg = 407 in. w.c.
Upon substitution, we obtain:
SP, = 0.034(118 - 5) (407) (1/[70 + 460] - 1/960) = 1.32 w.c.
10.4 Estimating Total Capital Investment
This section presents the information needed for estimating
the total capital investment (TCI) for hoods, ductwork, and
stacks The TCI includes the equipment cost (EC) for the hood,
ductwork or stack; taxes; freight charges; instrumentation (if.
applicable); and direct and installation costs. All costs are
presented in second quarter 1993 dollars, and are of "study"
estimate accuracy (±30 percent). Moreover, the costs are for
new facility installations; no retrofit costs are included.
The equipment costs are presented in Section 10.4.1, while
the installation costs are shown in Section 10.4.2. In each of
these sections, the three categories of equipment are covered in
separate subsections.
10.4.1 Equipment Costs
Several vendors provided costs (prices) for each of the
three equipment categories. Their responses reflected a range of
sizes, designs, and materials of construction. These prices have
been correlated against some easy-to-determine design (sizing)
parameter via least-squares regression analysis. Each of these
correlations pertains to a certain type of equipment (e.g.,
circular canopy hoods) within a specified size range of the
parameter in question (e.g., 2 to 200 ft2 inlet area). For that
reason a cost correlation should not be extrapolated outside the
10-41
-------
parameter range specified.
Some of the prices the vendors provided pertain to stock
("off-the-shelf") items, while other costs are for custom-
fabricated equipment. Vendors tend to specialize in either stock
or custom items. Most hoods and stacks are custom-made, either
fabricated in the vendor's factory or erected on-site.
Conversely, ductwork components usually are stock items, though
larger pieces have to be custom-made. (Of course, there are
exceptions to this.) Finally, all prices given in the following
section are "F.O.B. (free-on-board) vendor," meaning that they
include neither freight nor taxes.
10.4.1.1 Hood Costs
In all, four vendors provided prices for hoods.55 These
prices covered the following types of hoods:
B5" Canopy—circular
«®- Canopy—rectangular
K3" Push-pull
KS* Side-draft
i®" Back-draft (slotted)
Descriptions and design procedures for these hoods are given
in Sections 10.2.1 and 10.3.2, respectively. As explained in
Section 10.3.2, hood costs have been found to correlate well with
the hood inlet or face area (Af, ft2) . Furthermore, the
functional form that best fits the cost-face area correlation
(equation) is the "power function", or:
Ch = aA," (10.35)
where: Ch = hood cost ($)
a,b = equation regression parameters
The values of the equation parameters vary according to hood _
type and material of construction. These parameters are shown in
Table 10.5.
Illustration: What would be the cost of the electroplating tank
canopy hood sized for the illustration in Section 10.3.2.? Assume
that the hood is fabricated of FRP.
Solution: Recall that the face area (A.) calculated for that hood
was 98.5 ft2. Because this is a circular canopy hood, the equation
parameters from Table 10.5 are: a = 123 and b = 0.575. (Note that
10-42
-------
Table 10.5 Parameters for Hood Cost Equation8
=
Type of
Hood
Canopy -
circular
Canopy -
rectangular
Push-pull
Side-draft
Backdraf t
(slotted)
Backdraf t
(slotted)
Backdraft
(slotted)
Backdraft
(slotted)
Backdraft
(slotted)
==
Fabrication
Material
FRP1"
FRP
FRP
FRP
PVC*'*
PVCn
PP**
FRP
Galvanized
Steel
Equation E
a
123
294
595
476
303
789
645
928
688
' _ .
'arameter
b
0.575
0.505
0.318
0.332
1.43
0.503
0.714
0.516
0.687
Equation
Range
(A,, ft2)
2-200
— — — — •
2-200
2-200
2-200
0.6-2.0§§
1.1-2.1
1.1-2.1
1.1-2.1
0.5-1.3.
Based on data received from hood vendors. (See Reference
52.
f Fiberglass-reinforced plastic.
* Polyvinyl chloride.
• Each hood is equipped with two rows of slots, but no
dampers.
§
-------
this hood area falls within the equation range of 2 to 200_ft2.)
Substituting these parameters into equation 10.35, we obtain:
Ch = 123 (98. 5) a575 = $1,720.
10.4.1.2 Ductwork Costs
Several vendors provided ductwork prices, also for a range of
sizes, materials, and designs.56 These prices covered the
following equipment items:
I®1 Straight ductwork:
* Circular
A Steel sheet (galvanized carbon, w/ & w/o
insulation; 304 stainless;)
A Steel plate (coated carbon; 304 stainless)
A Plastic (FRP; PVC)
4 Square , . .
A Steel (aluminized carbon; w/ & w/o insulation)
«sr Elbows (90°) :
* Steel (galvanized carbon, w/ & w/o insulation;
304 stainless)
4 Plastic (FRP; PVC)
"3" Dampers:
4 Butterfly . .
A Steel (galvanized carbon, w/ & w/o insulation;
A Plastic (FRP; PVC, w/ & w/o actuators)
4 Louvered
A Steel (aluminized carbon w/ & w/o actuators)
4 Blast gate
A Steel (carbon)
A PVC
These prices were regressed against the diameter of the _
equipment i?em (straight duct, elbow, or damper) The regression
correlations were of three forms: power function (primarily),
exponential, and linear. Equation 10.35 depicts the power
function, while the other forms are:
10-44
-------
Exponential: C, = aebD (10.36)
Linear: C, = a + bD (10.37)
where: C, = cost of equipment item in question
a,b = regression parameters
The regression parameters are listed in Tables 10.6 to 10.-8,
along with the size applicability ranges for the respective
correlations. (Note: The correlations should not be extrapolated
outside these ranges.) The following paragraphs contain additional
information about the price data and the correlations:
f Straight duct: As indicated above, vendors provided prices for
steel plate, steel sheet (spiral-wound and longitudinal seam), and
plastic straight duct. The major difference between the two steel
duct types lies in the wall thickness. Steel plate duct typically
has wall thicknesses of from 3/16 in. to 1/2 in., while steel sheet
duct wall thicknesses usually range from 28 gauge to 10 gauge. As
Table 10.4 shows, this range corresponds to thicknesses of 0.0149
in to 0 1406 in., respectively, although the exact thicknesses
will vary with the type of steel used (e.g., carbon vs. stainless).
Also, as discussed in Section 10.3.3.2, each duct diameter can be
fabricated with a range of wall thicknesses.
Most of the steel duct vendors supplied prices for a minimum
and a maximum wall thickness for a given diameter. However, to_
simplify matters for cost estimators, these "low" and "high" prices
first were averaged, and then the average prices were regressed
against the diameters. This averaging was necessary, because those
making study cost estimates usually do not have enough information
available to predict duct wall thicknesses.
Prices for both circular and square insulated steel sheet duct
were among the data received. The insulated circular steel duct_is
"double-wall, spiral-wound" in construction, wherein the insulation
is installed between the inner and outer walls. Costs were
provided for both 1-in. and 3-in. fiberglass insulation
thicknesses. For the square duct, prices were given for a 4-in.
thickness of mineral wool insulation applied to the outer surface
of the duct. The correlation parameters in Table 10.6 reflect
these specifications.
Prices for both carbon steel (galvanized, painted, or
aluminized) and 304 stainless steel duct were received. The carbon
steel duct is used in situations where "mild" steel is suitable,
while the stainless steel duct is required whenever the gas stream
contains high concentrations of corrosive substances.
Vendors gave prices for plastic (FRP and PVC) duct_also (Table
10.8). However, for a given diameter this duct is fabricated in a
10-45
-------
Table 10.6 Parameters for Straight Steel Ductwork Cost Equations*
Duct
Type
Circular-
spiral
Circular-
spiral
Circular-
spiral
Circular-
spiral
Circular-
longitudinal"
Circular-
longitudinal
Circular-
longitudtnal
Circular-
longitudinal
Square
Square
— •
Material
Sheet -
galv CS*
Sheet -
304 SS*
Sheet -
galv CS
Sheet-
galv CS
Sheet -
galv CS
Sheet -
304 SS
Plate-
coat
csft
Plate-
304 SS**
Sheet-
alum
CS"
Sheet -
alum CS
Insulation
Thickness
(in.)
None
None
1
3
None
None
None
None
None
4
Equation
Type
Power
function
Power
function
Power
function
Power
function
Power
function
Power
function
Power
function
Power
function
Linear
Linear
Equation
Parameter
a
0.322
1.56
1.55
2.56
2.03
2.98
2.49
6.29
0.254
21.1
b
1.21
1 .00
0.936
0.937
0.784
0.930
1.15
1.23
2.21
5.81
Equation
Range
(D, in.)
3 - 84
3 - 84
3 - 82
3 - 82
6 - 84
6 - 84
6 - 84
6 - 84..
18 - 48
18 - 48
§ Based on data from ductwork vendors. (Reference 53.
f Spiral-wound and welded circular duct.
* Galvanized carbon steel sheet.
* 304 stainless steel sheet.
88 Circular duct welded along the longitudinal seam.
n Carbon steel plate with one coat of "shop paint".
** 304 stainless steel plate.
*• Aluminized carbon steel sheet.
10-46
-------
Table 10.7 Parameters for Steel Elbows and Dampers Cost Equations5
=====
Ductwork
Item
Elbows1'
Elbows
Elbows -
insulated§§
Dampers -
butterflyn
Dampers-
butterfly /insulated**
Damper s-
louvered"
Dampers-
louvered
w/actuatorsm
Dampers -
blast gates
=
Material
Galv CS*
304 SS*
Galv CS
Galv CS
Galv CS
Alum
cs«s
Alum CS
Carbon
steel
========
Equation
Type
Exponential
Exponential
Exponential
Exponential
Exponential
Power
function
Power
function
Power
function
=======
Equation
Parameter
a
30.4
74.2
53.4
23.0
45.5
78.4
208.
17.2
b
0.0594
0.0668
0.0633
0.0567
0.0597
0.860
0.791
0.825
====== —
Equation
Range
(D, in.)
6 - 84 "
6 - 60
3 - 78
4 - 40
4 - 40
18 - 48
18 - 48
3 - 18
53 .
§ Based on-data received from ductwork vendors. (See Reference
f Single-wall "gored" 90° elbows, uninsulated.
* Galvanized carbon steel sheet.
* 304 stainless steel sheet.
§§ Double-wall "gored" 90° elbows with 1-inch fiberglass
insulation.
ft Single-wall "opposed blade" type manual butterfly dampers.
** Double-wall "opposed blade" butterfly dampers with 1-inch
fiberglass insulation.
•* Louvered dampers with 95-98% sealing.
§§§ "Aluminized" carbon steel sheet.
nt Louvered dampers with electric actuators (automatic
controls).
10-47
-------
Table 10.8 Parameters for Plastic Ductwork Cost Equations^
Ductwork
Item
Straight
duct
Straight
duct
Elbows -90°
Elbows -90°
Dampers -
butterfly
Damper s-
butterf ly
Dampers-
butterfly w/actuators
Damper s-
blast gate
Material
PVCf
FRP*
PVC
FRP
PVC
FRP
PVC
PVC
Equation
Type
Power
function
Exponential
Power
function
Exponential
Power
function
Power
function
Exponential
Power
function
Equation
Parameter
a
0.547
11.8
3.02
34.9
10.6
35.9
299.
8.14
b
1.37
0.0542
1.49
0.0841
1.25
0.708
0.0439
1.10
-Equation
Range
(D, in.)
6 - 48
4 - 60
6 - 48
4 - 36
4 - 48
4 - 36
4 - 48 .
4 - 48
53
Based on data received from ductwork vendors. (See Reference
f Polyvinyl chloride.
* Fiberglass-reinforced plastic.
* Butterfly dampers with pneumatic actuators (automatic
controls). All other dampers listed in this table are manually-
controlled.
10-48
-------
single wall thickness, which varies from approximately 1/8 in. to
1/4 in. Consequently, the estimator is not required to select a
wall thickness when costing plastic duct.
1 Elbows: Prices for steel sheet and plastic 90° elbows were also
submitted. The steel sheet elbows were "gored" (sectioned) elbows
fabricated from five pieces of sheet metal welded together. Like
the straight duct, the steel elbows were priced at two wall
thicknesses: "minimum" and "maximum". These prices were averaged
before being regressed against the elbow diameter. Prices were
also given for both galvanized carbon steel elbows (with and
without 1-in. fiberglass insulation) and 304 stainless steel
elbows. Correlation parameters for steel elbows are listed in
Table 10.7.
Costs for both PVC and FRP 90° elbows were also given. The
PVC ells were fabricated from three sections ("three-piece
miter"), while the FRP elbows were one-piece molded fittings. As
with the plastic straight duct, each elbow of a given diameter was
fabricated in a single wall thickness. Table 10.8 contains
correlation parameters for plastic elbows.
\ Dampers: Prices were obtained for three types of dampers:
butterfly, louvered, and blast gates. The galvanized carbon steel
butterfly dampers were priced with and without 1-in. fiberglass
insulation, while prices for the aluminized carbon steel louvered
dampers were based on either manual or automatic control (via
electric actuators). Similarly, the PVC butterfly dampers were
manual or equipped with pneumatic actuators. Both the carbon steel
and the PVC blast gates were manual. Correlation parameters for
the steel and plastic dampers are shown in Tables 10.7 and 10.8, in
turn.
Illustration: A fabric filter handling 16,500 ft3/min of 200°F.
waste gas laden with noncorrosive cocoa dust is located 95 ft
across from and 20 ft above, the emission source (a drying oven).
Straight duct with four 90° elbows (all fabricated from spiral-
wound, galvanized carbon steel sheet) and a butterfly damper (also
galvanized CS) will be required to convey the gas from the source
to the control device. Assume that the ductwork is insulated to
prevent condensation. Estimate the cost of these items.
Solution: First, determine the diameter of the straight duct,
elbows, and damper. From Table 10.3, the minimum transport
velocity (ut) for cocoa dust is 3,000 ft/min. Substituting this
value and the gas volumetric flow rate into equation 10.27, we
obtain:
Dd = 1.128(16,500/3,000) 1/2 = 2.65 ft = 31.7 in.
Next, obtain the costs of the ductwork items as follows:
10-49
-------
us- straight ductwork: From Table 10.6, select the equation
parameters for galvanized circular spiral-wound duct (1-in.
insulation) and substitute them and the diameter into the
appropriate equation type (power function, equation 10.35).
Straight duct cost ($/ft) = 1. 55 (31. 7) a936 = $39.4/ft.
However, a total of 115 ft (95 + 20) of duct is required,, so:
Straight duct cost = $39.4/ft x 115 ft = $4,531.
es> Elbows: The Table 10.7 correlation parameters for galvanized
carbon steel, insulated elbows are 53.4 (a) and 0.. 0633 (b) .
However, the regression correlation form is exponential (equation
10.36). Thus:
Elbow cost ($) = 53.4e°-(*>33<31-7> = $397 ea.
For four elbows, the cost is: $397 x 4 = $1,588.
"^ Damper: Also from Table 10.7, select the correlation parameters
for galvanized carbon steel "dampers-butterfly/insulated" and
substitute into equation 10.36:
Damper cost ($) = 45. 5ea0597(3I'7) = $302.
After summing the above three costs, we obtain:
Total ductwork cost = $6,421 «= $6,420.
10.4.1.3 Stack Costs
Prices for steel and PVC short stacks were obtained from four
vendors.57 The steel stack costs were for those fabricated from
carbon and 304 stainless steels, both plate and sheet metal. As
with ductwork, the difference between steel sheet and plate lies in
the thickness. For these stacks, the sheet steel thickness ranged
from 18 to 16 gauge (0.05 to 0.06 in., approximately). Steel plate
thicknesses were considerably higher: 0.25 to 0.75 in, a fact that
makes them more resistant to wind and other loadings than stacks
fabricated of steel sheet. This is especially true for taller
stacks. The major drawback is that plate steel stacks are more
costly than those fabricated from steel sheet.
Another feature that increases costs is insulation. As the
correlation parameters show (Table 10.9), insulated stacks cost as
much as three times more per foot than uninsulated. With or
without insulation, a typical short (15-ft) steel stack consists of
r o
the following components:
10-50
-------
B3T Longitudinal seam duct (12-ft section)
BS- Reducer fitting (3-ft)
BS" Drip pan
Bar Support plate (1/4-in, welded to stack)
Bar Rectangular tap (for connecting to fan discharge)
«sr Ring (for attaching guy wires)
Taller stacks may require additional components, such as ladders
and platforms, guy wires or other supports, and aircraft warning
lights. (See Section 10.2.3.)
Table 10 9 lists the parameters and applicable ranges of the
stack cost correlations. The correlations cover short PVC stacks,
and taller stacks fabricated from plate steel (carbon and 304
stainless types) and sheet steel (insulated and uninsulated)
Except for three double-wall sheet steel designs, these stacks are
of single-wall construction.
Note that all of the correlations are power functions._ Also
note that the equations apply to various ranges of stack_height
In all but one of these equations the cost is expressed in $/ft of
stack height. The exception is the cost equation for insulated
carbon steel sheet stacks of heights ranging from 30 to 75 feet.
In this equation the cost is expressed in $.
This last cost equation is different in another respect The
other six equations in Table 10.9 correlate stack cost ($/ft) with
stack diameter (D., in.). However, this seventh equation correlates
stack cost with stack surface area (Ss/ ft2), a variable that
incorporates both the stack diameter and the stack_height (HJ . The
surface area is calculated via the following equation:
S, = (7r/12)DsHs (10.38)
where: 1/12 = stack diameter (D,) conversion factor
Illustration: Estimate the cost of the stack sized in the Section
10.3.4.3 illustration.
Solution: Recall that the stack dimensions were: Hs = 95 ft and
D = 2 12 ft = 25 4 in. Both dimensions fall within the ranges of
the co'st correlations for steel plate stacks. Because the previous
illustration did not indicate whether the waste gas was corrosive,
we will estimate the prices for both carbon steel and 304 stainless
10-51
-------
Table 10.9 Parameters for Stack Cost Equations'
I
Material
PVC§§
Plate coated CS~
Plate-304 SSW
Sheet -qalv CS"
Sheet- 304 SS§i>§
Sheet- insul CS/DWnt
Sheet -uninsul CS/DW***
Sheet -insul
CS/DW"*
• •
Equation I
a
0.393
3.74
. —
12.0
2.41
4.90
_ —
143.
10.0
142.
••
3arametert
b
1.61
1.16
1.20
1.15
1.18
0.402
1.03
0.794
Equatio
Ds (in)*
~
12 - 36
6 - 84
6 - 84
8 - 36
8 - 36
18 - 48
18 - 48
24 - 48
n Range
Hs (ft)*
< 10
,^____ — .,. , _ .,_.
20 - 100
20 - 100
< 75
< 75
< 15
< 15
30 - 75
« Based on data received from stack vendors. (See Reference
54. )
* All cost equations are power functions. (See equation
10.35.) Except where noted, costs are expressed xn terms of $/ft
of stack height.
* Stack diameter range to which each equation applies.
• Stack height range to which each equation applies.
§§ Polyvinyl chloride.
tt carbon steel plate with one coat of "shop paint".
** 304 stainless steel plate.
•* Galvanized carbon steel sheet.
§§§ 304 stainless steel sheet.
™ Aluminized carbon steel sheet covered with 4 inches of
fiberglass insulation (double-wall construction).
** Uninsulated aluminized carbon steel sheet (double-wall
construction).
•- Costs for these stacks are expressed in $, and are
correlated with the stack surface area (Sif ft ) .
10-52
-------
steel plate stacks.
Upon substituting the equation parameters and stack dimensions
into equation 10.35, we obtain:
Price (carbon steel) = 3 . 74 (25 .4) h16 ($/ft) x 95 ft
= $15,100.
Price (304 stainless) = 12 . 0 (25 .4) L2° ($/ft) x 95 ft
= $55,300.
Notice that the price of the stainless steel stack is nearly
four times that of the carbon steel stack. In view of this
difference the estimator needs to obtain more information on the
waste gas stream properties, so that he/she can select the most
suitable stack fabrication material. Clearly, it would be_a poor
use of funds to install a stainless steel stack where one is not
needed.
10.4.2 Taxes, Freight, and Instrumentation Costs
Taxes (sales, etc.) and freight charges apply to hoods,
ductwork, and stacks, as they do to the control devices that these
auxiliaries support. As discussed in Chapter 2, these costs vary,
respectively, according to the location of the ventilation system
and the site's distance from the vendor. Typical values are 3*
(taxes) and 5% (freight) of the total equipment cost.
Unlike the control devices, ventilation systems generally are
not instrumented. The exception would be an electric or pneumatic
actuator for a butterfly or louvered damper. In such a case
however the cost of the instrument (actuator and auxiliaries)
would be included in the damper price. Thus, no supplementary
instrumentation cost is included.
10.4.3 Purchased Equipment Cost
With ventilation systems, the purchased equipment cost (PEC,)
is the sum of the equipment, taxes, and freight costs.
Incorporating the typical values listed in Section 10.4.2, we
obtain:
PECt = EC, + 0.03EC, + 0.05EC,
= 1.08(ECt) (10-39)
where: EC, = total cost of hood(s), ductwork, and stack(s)
10--53
-------
10.4.4 Installation Costs
When making a cost estimate for an air pollution control
svstem according to the procedure in this manual, the estimator
first determines the cos? of the control device, then estimates the
costs of -such auxiliaries as the hood, ductwork stack fan and
motor and other items. To these items he/she adds the costs of
TSs?rumSn?ation, taxes, and freight, to obtain the ^ * ally,
the estimator multiplies the PEC by the insDilation ff£tor
appropriate to the control device e.g 2 20 for gas ab orbers)
3 to 9 for more information about these factors.)
For this reason, it usually is unnecessary to estimate the
installation cost of the ventilation system separately. However
there" may Te occas?ons where a hood, a stack, or ductwork has to be
Called alone either as replacement equipment or to augment the
existing venation system. In those instances, the estimator may
want to estimate the cost of installing this item.
As might be imagined, these installation costs vary
considerably, according to geographic location *%**££*<>"* °f
the facility, equipment design, and sundry other variables.
Nonetheless some of the vendors (and a peer reviewer59) provided
razors for'hoods and ductwork, which, when multiplied by their
respective purchased equipment costs, will yield approximate
installation costs. These are:
B3" Hoods: 50 to 100%
us" Ductwork: 25 to 50%
If one or both of the latter factors is used, the total
capital investment (TCI) of the hood and/or ductwork would be:
TCI = (1 + IFh/d) x PECh/d (10.40)
where- IFh/d = installation factor for hood(h)/ductwork(d) .
PEcCa = purchased equipment cost of hood(h)/ductwork(d)
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10.5 Estimating Total Annual Cost
10.5.1 Direct Annual Costs
electricity cost (Cc, $/yr) can be calculated as follows.
Cc = (1.175 x 10^)PeQFd0/e (10.41)
where- p = electricity price ($/kwh)
wnere. p = ^^ ^^ ^^ (actual ft3/min)
n = waste gda j-j-i_iw j-cn-^-. \^.^~ — ,
F = static pressure drop through ventilation
system (in. w.c.)
6 = operating factor (hr/yr)
e = combined fan-motor efficiency
..
consumed by the fan needed to convey the gas
of 0.6,
^J_ J_ dii lll\_/ V-V^J- — — - - -t
factor.
Solution: . Recall that the P-ssure^rop^nd^a^flow r^e for this
respectively.^Upon"substituting these values and the other
parameters into equation 10.40, we obtain:
Ce = (1-175 x 10^) (0.075) (15,000) (0.313) (8,000)/0.6
= $552/yr.
'3 Technically, this direct annual cost should be allocated to
equation has beln'provided as a temporary convenience to Manual
users
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10.5.2 Indirect Annual Costs
The indirect annual costs for ventilation systems include
property taxes, insurance, general and administrative (G&A), and
capital recovery costs. (Overhead—a fifth indirect annual
cost—is not considered, because it is factored from the sum of_the
operating, supervisory, maintenance labor and maintenance materials
costs, which is negligible.) When a ventilation system is part of
a control system, these costs are included in the control system -
indirect annual cost. However, if the ventilation equipment has
been sized and costed separately, these costs can be computed from
the total capital investment (TCI) via standard factors, as
follows:
Indirect Annual Cost
Computation Equation
Property taxes_
0.01 x TCI
Insurance
0.01 x TCI
General and Administrative
0.02 x TCI
Capital recovery
CRF x TCI
The "CRF" term in the capital recovery equation is the capital
recovery factor, which is a function of the economic life of the
ventilation system and the interest rate charged to the total
capital investment. (See Section 2.3 of this manual for more
discussion of the CRF and the formula used for computing it.)
For a ventilation system, the economic life varies from at
least 5 to 10 years to 15 to 20 years or more.60'61 In general, the
ventilation equipment should last as long as the control system it
supports As discussed in Section 2.3, the interest rate to use in
the CRF computation should be a "pre-tax, marginal (real) rate of
return" that is appropriate for the investor. However, for those
cost analyses related to governmental regulations, an appropriate
"social" interest (discount) rate should be used. For these kinds
of analyses the Office of Management and Budget (OMB) directs that
a real annual interest rate of 7% be used.62 (This replaces the
10% rate OMB previously had mandated.)
10.5.3 Total Annual Cost
The total annual cost (TAG) is calculated by adding the direct
(DC) and indirect (1C) annual costs:
TAG = DC + 1C
[10.42
10-56
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10.6 Acknowledgements
Several firms and individuals provided very useful technical
and cost information to this chapter. Foremost among these was
Todd N. Stine of United McGill Corporation (Raleigh NC), who
submitted current prices for a variety of ductwork items, as well
as a comprehensive product catalog and engineering design manual.
in addition, Mr. Stine patiently replied to the author's many
Questions, providing supplemental data when requested. Samir
Srandikar of EPCON Industrial Systems (The Woodlands TX) and
Gregory P Michaels of Piping Technology & Products (Houston, TX)
also were very helpful in submitting data and responding to
inquiries.
The author also would like to thank the following firms for
their valuable contributions:
03" Air Plastics, Inc. (Mason, OH)
BS> General Resource Corporation (Hopkins, MN)
ear Harrington Industrial Plastics, Inc. (Chino, CA)
es- Intellect Systems & Marketing, Inc. (Bohemia, NY)
us- Wer-Coy Metal Fabrication Co. (Warren, MI)
in addition, several individuals reviewed the draft chapter
and provided valuable suggestions, supplemental information, or
both The EPA peer reviewers, all located at Research Triangle
Park, NC, were:
B3> James C. Berry (OAQPS/ESD)
03- Peter A. Eckhoff (OAQPS/TSD)
cr Norman Kaplan (ORD/AEERL)
car James H. Maysilles (OAQPS/ESD)
•5- Larry Sorrels (OAQPS/ESD)
Finally Howard Goodfellow of Goodfellow Consultants, Inc.
(Mississauga, Ontario, Canada) also reviewed the chapter and
supplied helpful comments.
10-57
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References
1 Goodfellow, H.D. "Ancillary Equipment for Local Exhaust
v;ntilation Systems" In: Air Pollution Engineering Manual. New
York: Van SosSanS Reinhold/Air and Waste Management Association.
1992, pp. 155-205.
2. Burton, D. Jeff. Industrial Ventilation Work Book. Salt Lake
City: DJBA, Inc. 1989.
3 The Measurement Solution: Using a Temporary Total Closure
f'nr- Canture Efficiency Testing. Research Triangle Park, NC: U.S.
E?viro^en?af ProtScJon Agency. August 1991 (EPA-450/4-91-020),
pp. 3,11.
4. The Measurement Solution, pp. 11-29.
5 Heinsohn, Robert Jennings. Industrial Ventilation:
Engineering Principles. New York: John Wiley & Sons, Inc. 1991.
6 Telephone conversation between William M. Vatavuk, U.S.
Environmental Protection Agency (Research Triangle Park NC) and
Todd N. Stine, United McGill Company (Raleigh, NC), May 24, 1993.
7 Thermoplastic Duct (PVC) Construction Manual. Vienna, VA:
Sheet Metal ard Air Conditioning Contractors' National
Association, Inc. (SMACNA). May 1987, pp. 61-85.
8. Thermoplastic Duct Construction Manual, p. 64.
9. Burton, p. 6-7.
10. Dust Control System Accessories Price List. Huntington Park,
CA: Murphy-Rodgers, Inc. July 1992.
11 Price and Data Catalog: Standard Ductwork Components.
Warren, MI: Wer-Coy Metal Fabrication Co. 1992-93.
12 Letters from Samir Karandikar, EPCON Industrial Systems
(Woodlands TX) to William M. Vatavuk, U.S. Environmental
ProScSon Agency (Research Triangle Park, NC) . May 21 and June
9, 1993.
13 -Double Wall Insulated Duct and Fittings." In: Sheet Metal
Division Catalog. Groveport, OH: United McGill Corporation.
1990
10-58
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14 "Single-Wall Round and Flat Oval Duct and Fittings." In:
Sheet Metal Division Catalog. Groveport, OH: United McGill
Corporation. 1990.
15 HVAC Duct Construction Standards: Metal and Flexible.
Vienna VA- Sheet Metal and Air Conditioning Contractors'
National Association, Inc. (SMACNA). 1985, pp. 2-15 to 2-17.
16 Wherry, T.C. and Peebles, Jerry R., "Process Control". In:
Perry's Chemical Engineers' Handbook, Sixth Edition. New York:
McGraw-Hill, Inc. 1984.
17. Product catalog. Rio, WI: Gaskets, Inc. 1994.
18. HVAC Duct Construction Standards, pp. 4-2 to 4-3.
19. HVAC Duct Construction Standards, pp. 4-2 to 4-7.
20 Letter from Howard D. Goodfellow, Goodfellow Consultants
(Mississauga, Ontario, Canada) to William M. ^Vatavuk, U.S
Environmental Protection Agency (Research Triangle Park, NC) .
February 23, 1994.
21 Guide for Steel Stack Design and Construction. Vienna, VA
Sheet Metal and Air Conditioning Contractors' National
Association, Inc. (SMACNA). 1983.
22. Goodfellow, pp. 192-193.
23. Goodfellow, p. 193.
24 Peters Max S. and Timmerhaus, Klaus D. Plant Design and
Economics for Chemical Engineers, Third Edition. New York:
McGraw-Hill, Inc., 1980, pp. 508-510.
25. Burton, pp. 2-10 to 2-11.
26. Burton, p. 2-11.
27. Burton, pp. 4-5 to 4-8.
28. Burton, p. 5-12.
29. Burton, pp. 5-15 to 5-16.
30. Burton, p. 5-5.
31. Burton, pp. G-2, G-5.
32. Burton, p. 5-18.
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33. Vatavuk, William M. and Neveril, Robert B "Estimating Costs
of Air-Pollution Control Systems, Part III: Est^ing the Size
and Cost of Pollutant Capture Hoods," Chemical Engineering,
December 1, 1980, pp. Ill to 115.
34. Telephone conversation between William M Vatavuk U.S
Environmental Protection Agency (Research Triangle Park NC) and
Dennis Woll, Air Plastics, Inc. (Mason, OH), August 10, 1993.
35. Telephone conversation between William M Vatavuk U.S
Environmental Protection Agency (Research Triangle Park NC) and
Pat Caputo, intellect Systems & Marketing, Inc. (Bohemia, NY),
October 22, 1993.
36. Burton, "Chart 9".
37 Letter from Todd N. Stine, United McGill Corporation
(Raleigh, NC) to William M. Vatavuk, U.S. Environmental
PrStec?ion Agency (Research Triangle Park, NC). June 10, 1993.
38. Green, Don W. and Maloney, James 0. Perry's Chemical
Engineers' Handbook, Sixth Edition. New York: McGraw-Hill, Inc.
1984.
39 Peters, Max S. and Timmerhaus, Klaus D. Plant Design and
Economics for Chemical Engineers, Fourth Edition. New York:
McGraw-Hill, Inc. 1991.
40. Engineering Design Reference Manual for Supply Air Handling
Systems. Groveport, OH: United McGill Corporation. 1992, pp. 3-
4.
41. Engineering Design Reference Manual, p. 8.
42. Burton, "Chart 5".
43. Engineering Design Reference Manual, p. 7.
44. Burton, p. 6-6.
45. Burton, "Chart 13".
46. Goodfellow, p. 193.
47. Guide for Steel Stack Design and Construction, pp. 39 to 50.
48. Goodfellow, p. 193.
49. Carlton-Jones, Dennis and-Schneider, H.B., "Tall Chimneys,"
Chemical Engineering, October 14, 1968, p. 167.
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50. Guideline for Determination of Good Engineering Practice:
Stack Height (Technical Support Document for Stack Height
Regulations) (Revised). Research Triangle Park, NC: U.S.
Environmental Protection Agency. June 1985 (NTIS PB-85-225241),
p.l.
51. Guideline for Determination of Good Engineering Practice, pp.
50-51.
52. Guideline for Determination of Good Engineering Practice, pp.
1-2.
53. Goodfellow, p. 194.
54. Guide for Steel Stack Design and Construction, p. 4.
55. Hood cost data request responses from four hood vendors to
William M. Vatavuk, U.S. Environmental Protection Agency
(Research Triangle Park, NC). June-July 1993.
56 Ductwork cost data request responses from six vendors to
William M. Vatavuk, U.S. Environmental Protection Agency
(Research Triangle Park, NC). May-July 1993.
57. Stack cost data request responses from four vendors to
William M. Vatavuk, U.S. Environmental Protection Agency
(Research Triangle Park, NC). May-July 1993.
58. Op. cit., Stine-Vatavuk letter, June 10, 1993..
59. Goodfellow-Vatavuk letter.
60. Goodfellow-Vatavuk letter.
61 Telephone conversation between William M. Vatavuk, U.S.
Environmental Protection Agency (Research Triangle Park, NC) and
Todd N. Stine, United McGill Company (Raleigh, NC), December 10,
1993 .
62 Darman, Richard. Guidelines and Discount Rates for Benefit-
Cost Analysis of Federal Programs (OMB Circular No. A-94,
Revised). Washington, DC: Office of Management and Budget.
October 29, 1992.
10-61
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