United States
Environmental Protection
Agency
Office of Air Quality
Planning and Standards
Research Triangle Park NC 27711
November 1992
EPA-453/D-92-016b
Air
&EPA Hazardous Air Pollutant Draft
Emissions from Process Units EIS
in the
Synthetic Organic Chemical
Manufacturing Industry--
Background Information
for Proposed Standards
Volume 1B: Control Technologies
-------
EPA-453/D-92-016b
Hazardous Air Pollutant Emissions
from Process Units in the
Synthetic Organic Chemical
Manufacturing Industry--
Background Information
for Proposed Standards
Volume IB: Control Technologies
Emission Standards Division
U.S. Environmental Protection Agency
Region 5, Library (Pi-T")
77 West Jackson BcuU^/d, 12th Floor
Chicago, IL 60604-3590
U.S. Environmental Protection Agency
Office of Air and Radiation
Office of Air Quality Planning and Standards
Research Triangle Park, North Carolina 27711
November 1992
-------
(DISCLAIMER)
This Report has been reviewed by the Emission Standards Division of
the Office of Air Quality Planning and Standards, EPA, and approved
for publication. 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, N.C. 27711, or from the National Technical
Information Service, 5285 Port Royal Road, Springfield, VA 22161.
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ENVIRONMENTAL PROTECTION AGENCY
Background Information for Proposed Standards
Hazardous Air Pollutant Emissions from Process Units
in the Synthetic Organic Chemical Manufacturing Industry
Volume IB: Control Technologies
Prepared by:
_
Bruce Jordan (Date)
Directory Emission Standards Division
U.S. Environmental Protection Agency
Research Triangle Park, N.C. 27711
1. The proposed standards would regulate emissions of organic
hazardous air pollutants (HAP's) emitted from chemical
manufacturing processes of the Synthetic Organic Chemical
Manufacturing Industry (SOCMI). Only those chemical
manufacturing processes that are part of major sources under
Section 112 (d) of the CAA would be regulated. The recommended
standards would reduce emissions of 149 of the organic
chemicals identified in the CAA list of 189 hazardous air
pollutants.
2. Copies of this document have been sent to the following
Federal Departments: Labor, Health and Human Services,
Defense, Office of Management and Budget, Transportation,
Agriculture, Commerce, Interior, and Energy; the National
Science Foundation; and the Council on Environmental Quality.
Copies have also been sent to members of the State and
Territorial Air Pollution Program Administrators; the
Association of Local Air Pollution Control Officials; EPA
Regional Administrators; and other interested parties.
3. The comment period for this document is 90 days from the date
of publication of the proposed standard in the Federal
Register. Ms. Julia Stevens may be contacted at 919-541-5578
regarding the date of the comment period.
4. For additional information contact:
Dr. Janet Meyer
Standards Development Branch (MD-13)
U.S. Environmental Protection Agency
Research Triangle Park, N.C. 27711
Telephone: 919-541-5299
5. Copies of this document may be obtained from:
U.S. EPA Library (MD-35)
Research Triangle Park, N.C. 27711
-------
National Technical Information Service (NTIS)
5285 Port Royal Road
Springfield, VA 22161
Telephone: 703-487-4650
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OVERVIEW
Emission standards under Section 112(d) of the Clean Air Act
apply to new and existing sources in each listed category of
hazardous air pollutant emission sources. This background
information document (BID) provides technical information used in
the development of the Hazardous Organic National Emission
Standard for Hazardous Air Pollutants (NESHAP), which will affect
the Synthetic Organic Chemical Manufacturing Industry (SOCMI).
The BID consists of three volumes: Volume 1A, National Impacts
Assessment (EPA-453/D-92-016a); Volume IB, Control Technologies
(EPA-453/D-92-016b); and Volume 1C, Model Emission Sources
(EPA-453/D-92-016C).
Volume 1A presents a description of the affected industry
and the five kinds of emission points included in the impacts
analysis: process vents, transfer loading operations, equipment
leaks, storage tanks, and wastewater collection and treatment
operations. Volume 1A also describes the methodology for
estimating nationwide emissions, emission reductions, control
costs, other environmental impacts, and increases in energy
usage resulting from a potential NESHAP; and presents three
illustrative sets of potential national impacts and a summary of
the economic analysis. While Volume 1A provides the overview of
how information on model emission sources and control technology
cost were used to estimate national impacts, Volumes IB and 1C
contain detailed information on the estimation of control
technology performance and costs and model emission source
development.
Volume IB discusses the applicability, performance, and
costs of combustion devices; collection systems and recovery
devices; storage tank improvements; and control techniques for
equipment leak emissions. These control technologies were the
basis of the Hazardous Organic NESHAP impacts analysis. These
control technologies are applicable to emission points in the
SOCMI and in other source categories. Methods for estimating
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capital costs and annual costs (including operation and
maintenance costs) of each control technology are presented.
Volume 1C presents descriptions of each kind of emission
point included in the impacts analysis and the development of
model emission sources to represent each kind of emission point
for use in the impacts analysis. The emission reductions, other
environmental impacts, and energy impacts associated with
application of the control technologies described in Volume IB to
the model emission sources is discussed. For illustrative
purposes, the environmental, energy, and cost impacts that would
results from control of several example model emission sources
are presented.
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TABLE OF CONTENTS
Section Page
LIST OF TABLES ix
LIST OF FIGURES xi
ACRONYM AND ABBREVIATION LIST xii
1.0 INTRODUCTION 1-1
2.0 EMISSION CONTROL TECHNOLOGIES 2-1
2.1 Combustion Control Devices 2-1
2.1.1 Flares ..... 2-1
2.1.2 Thermal Incinerators ......... 2-7
2.1.3 Catalytic Incinerators 2-13
2.1.4 Industrial Boilers and Process Heaters 2-16
2.2 Collection Systems and Recovery Devices . . . 2-20
2.2.1 Vapor Collection Systems for
Loading Racks 2-21
2.2.2 Condensers 2-27
2.2.3 Steam Strippers 2-33
2.2.4 Carbon Adsorbers 2-43
2.2.5 Absorbers 2-48
2.3 Storage Tank Improvements for Emission
Reduction 2-52
2.3.1 Description of Tank Improvements . . . 2-53
2.3.2 Factors Affecting Control Efficiency . 2-54
2.3.3 Applicability of Storage Tank
Improvements 2-62
2.4 Equipment Leak Emission Sources and Emission
Control Techniques ... 2-63
2.4.1 Equipment Description and
Specifications .'.....,..... 2-65
2.4.2 Closed Vent Systems and Control
Devices . . . . . . . . 2-77
2.4.3 Work Practices 2-78
2.5 References . 2-84
3.0 COST ANALYSIS 3-1
3.1 Cost Methodology for Combustion Systems . . . 3-1
VII
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TABLE OF CONTENTS
(Continued)
Section
3.1.1 Cost Methodology for Flare Systems . . 3-1
3.1.2 Cost Methodology for Incinerator
Systems 3-7
3.2 Cost Methodology for Collection Systems and
Recovery Devices 3-21
3.2.1 Cost Methodology for Vapor Collection
Systems for Loading Racks 3-21
3.2.2 Cost Methodology for Condensers ... 3-24
3.2.3 Cost Methodology for Steam Stripping . 3-31
3.3 Cost Methodology for Storage Tank
Improvements 3-40
3.3.1 Design Considerations Affecting Cost . 3-40
3.3.2 Development of Capital Costs 3-43
3.3.3 Development of Total Annual Costs . . 3-44
3.4 Cost Methodology for Equipment Leak Control
Technologies 3-45
3.4.1 Control Equipment 3-45
3.4.2 Leak Detection and Repair Techniques . 3-50
3.4.3 Capital Costs 3-57
3.4.4 Annual Costs 3-57
3.5 References 3-61
Vlll
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LIST OF TABLES
Table Page
2-1 Controlled and Uncontrolled Internal
Floating Roof Deck Fittings 2-55
2-2 Effectiveness of Internal Floating Roofs
on an Example Tank 2-57
2-3 Internal Floating Roof Rim Seal Systems,
Seal Loss Factors, and Control Efficiencies . 2-60
3-1 Flare General Design Specifications 3-2
3-2 Bases and Factors for Annual Costs for Flares 3-6
3-3 Incinerator General Design Specifications . . 3-8
3-4 Scrubber General Design Specifications . . . 3-12
3-5 Capital Cost Factors for Thermal Incinerators 3-18
3-6 Bases and Factors for Annual Costs for
Thermal Incinerators 3-20
3-7 Total Annual Cost for Loading Rack
Vapor Collection Systems 3-25
3-8 Equipment Cost Equations for Packaged
Refrigerated Condenser Systems 3-27
3-9 Bases and Factors for Annual Costs for
Refrigerated Condenser Systems 3-29
3-10 Equipment Cost Equations for a Steam
Stripping Unit 3-33
3-11 Cost Methodology for Estimating Total Capital
Investment for a Steam Stripping System ... 3-36
3-12 Cost Methodology for Estimation of Total
Annual Cost for a Steam Stripping System . . 3-41
3-13 Base Cost Data for Equipment Leak
Control Devices 3-47
3-14 Base Costs and Assumptions for a
Leak Detection Program 3-54
IX
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LIST OF TABLES
(Continued)
Table Page
3-15 Equations for Determining Costs and Number of
Leaks for a Leak Detection Program 3-55
3-16 Derivation of Annualized Costs for Control
of Equipment Leaks 3-58
x
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LIST OF FIGURES
Figure Page
2-1 Steam-Assisted Elevated Flare System .... 2-3
2-2 Discrete Burner, Thermal Incinerator .... 2-9
2-3 Distributed Burner, Thermal Incinerator . . . 2-11
2-4 Catalytic Incinerator 2-15
2-5 Vapor Balancing System 2-24
2-6 Refrigerated Condenser System for
VOC Vapor Recovery 2-28
2-7 Schematic Diagram of a Shell and Tube Surface
Condenser 2-30
2-8 Continuous Steam Stripper System 2-36
2-9 Two-Stage Regenerative Adsorption System . . 2-45
2-10 Packed Tower Absorption Process 2-50
3-1 Summary of Total Capital Investment
versus Wastewater Feed Rate for Steam
Stripping Unit 3-37
3-2 Unit Operating Costs Versus Wastewater
Feed Rate for Steam Stripping Unit 3-42
XI
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ACRONYM AND ABBREVIATION LIST
API
Btu
cfm
cm
CO 2
EPA
fpm
ft
gal
gm
gin-mole
H2
HAP
HON
hp
hr
in.
kg
kPa
t
Ib
Ib-mole
LEL
MACT
Mg
mm
MMBtu
MJ
MW
N2
NESHAP
American Petroleum Insititute
British thermal unit(s)
cubic foot (feet) per minute
centimeter
carbon dioxide
U.S. Environmental Protection Agency
foot (feet) per minute
foot (feet)
gallon(s)
gram(s)
gram-mole(s)
hydrogen
hazardous air pollutant
hazardous organic NESHAP
horsepower
hour(s)
inch(es)
kilogram(s)
kilopascal
liter(s)
pound(s)
pound mole(s)
lower explosive limit
maximum achievable control technology
megagram(s)
millimeter(s)
million British thermal unit(s)
megajoule(s)
megawatt(s)
nitrogen
National Emission Standards for Hazardous
Air Pollutants
XII
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NPDES
NSPS
OAQPS
OCCM
OSHA
POTW
ppmv
psia
RACT
RD/PRV
scf
scfm
scm
sec
SOCMI
TLV
VOC
VOL
tiin
W
W/m2
yr
$
National Pollutant Discharge Elimination
System
New Source Performance Standard(s)
Office of Air Quality Planning and Standards
OAQPS Control Cost Manual
Occupational Safety and Health
Adiminstration
publicly owned treatment works
part(s) per million by volume
pound(s) per square inch absolute
reasonably available control technology
rupture disk/pressure relief valve
standard cubic foot (feet)
standard cubic foot (feet) per minute
standard cubic meter
second(s)
synthetic organic chemical manufacturing
industry
threshold limit value
volatile organic compound
volatile organic liquid
micrometer(s)
watt(s)
watts per square meter
year(s)
dollar(s)
Xlll
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1.0 INTRODUCTION
The hazardous organic national emission standards for
hazardous air pollutants (NESHAP) will address the following
five sources of emissions:
• Equipment leaks;
• Transfer operations;
• Process vents;
• Storage vessels (tanks); and
• Wastewater treatment operations.
This volume of the background information document (BID)
for the hazardous organic NESHAP (HON) presents the
technologies considered for control of organic hazardous air
pollutant (HAP) emissions from these sources.
The control technologies described here are common in the
synthetic organic chemical manufacturing industry (SOCMI) and
have been studied by the U.S. Environmental Protection Agency
(EPA) during past rulemaking. The discussions of efficiency,
applicability, and costs presented in this volume are a
compilation of information gathered during those earlier
studies.
The emission sources affected by the HON have the
potential to emit organic HAP's, which are part of the larger
criteria pollutant class of volatile organic compounds
(VOC's). The control technologies discussed in this volume
reduce emissions of VOC's and therefore organic HAP's.
Because data on the performance of the control technologies
are based on reduction of VOC's, these discussions present
reductions in terms of VOC's. Similar emission reduction
efficiencies would be expected for the smaller group of
organic HAP's.
1-1
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For each of the control technologies discussed in
Chapter 2, the design and operation of the control device or
system is described, including an explanation of the physical
or chemical processes that destroy the VOC or remove it from
the emission stream. The second part of each discussion
describes the factors affecting the efficiency of the control
device, including vent stream characteristics and control
device operating parameters. The last part of each discussion
addresses the applicability of the technologies; for example,
to what types of SOCMI sources can the technology be
successfully applied?
Most of the control technologies discussed in Chapter 2
may be applied to many different types of emission sources.
For this reason, the discussions of the technologies were
intended to be independent of emission source. However,
because vapor collection systems for loading racks, storage
tank improvements, and equipment leak controls are specific to
an emission source, the discussions of these technologies also
include a description of the emission source.
Chapter 2 contains descriptions of the following
technologies:
• Combustion devices (flares, thermal and catalytic
incinerators, and industrial boilers and process
heaters);
• Collection systems and recovery devices (vapor
collection systems for loading racks, condensers,
steam strippers, carbon adsorbers, and absorbers);
• Storage tank improvements; and
• Control techniques for equipment leak emissions.
For combustion devices and collection and recovery
systems, descriptions of the technologies include a discussion
of the factors affecting their efficiency and their
applicability to different waste streams.
For storage tanks, the discussion focuses on installing a
floating roof and upgrading an existing floating roof. The
types and mechanisms of different VOC losses are described
with the controls applicable to the particular loss types.
1-2
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The discussion of equipment leak emissions describes the
various sources of equipment leak emissions, that is, the
mechanisms by which different components become emitters of
VOC's. Alternative (leakless) components, closed vent
systems, and leak detection and repair (LDAR) programs are
discussed as methods for reducing equipment leak emissions.
Chapter 3 explains the procedures for calculating the
cost of applying these different emission reduction
technologies. The cost methodology includes discussion of
design considerations affecting costs, development of capital
costs, and development of annual costs. Of the combustion
technologies, flares and incinerator systems (including an
acid gas scrubber) are addressed. Cost methodologies for
collection systems and recovery devices are presented for
loading rack vapor collection systems, condensers, and steam
strippers. For storage tanks, cost methodologies are
described for installing a new floating roof and upgrading an
existing floating roof. Finally, for equipment leaks, cost
methodologies are included for the control equipment
applicable to certain source types and for LDAR programs.
The appendices to this volume contain example cost
calculations for applying control technologies to the five
emission sources covered by the HON.
1-3
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2.0 EMISSION CONTROL TECHNOLOGIES
This chapter discusses technologies for control of
organic HAP's from sources in the SOCMI. Because organic
HAP's are VOC's, technologies that have been used to control
VOC emissions were evaluated for control of organic HAP's.
These technologies include combustion devices, collection and
recovery devices, improvements to reduce emissions from
storage tanks, and equipment and work practices that reduce
emissions from equipment leaks.
2.1 COMBUSTION CONTROL DEVICES
This section discusses devices that control emissions of
VOC's by means of combustion. Combustion control devices,
unlike noncombustion control devices, alter the chemical
structure of the VOC. Destruction of the VOC through
combustion is complete if all VOC's are converted to carbon
dioxide and water. Incomplete combustion results in some of
the VOC remaining unaltered or being converted to other
organic compounds such as aldehydes or acids.
The combustion control devices discussed in the following
four subsections include flares, thermal incinerators,
catalytic incinerators, and boilers and process heaters. The
discussion for each device treats its typical design and
operation, factors affecting destruction efficiency, and
applicability.
2.1.1 Flares
2.1.1.1 Description of Flares. Flaring is an open
combustion process in which the oxygen necessary for
combustion is provided by the air around the flame. Good
combustion in a flare is governed by flame temperature,
residence time of organic species in the combustion zone,
turbulent mixing of the organic species to complete the
2-1
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oxidation reaction, and the amount of oxygen available for
free radical formation.
Many flare systems are operated with baseload gas
recovery systems to recover VOC's from the flare header system
for reuse. The recovered VOC's may be used as feedstock in
other processes or as a fuel in process heaters, boilers, or
other combustion devices. When baseload gas recovery systems
are applied, the flare is typically used to combust process
upset and emergency gas releases that the baseload system is
not designed to recover. The operation of a baseload gas
recovery system may offer an economic advantage over operation
of a flare alone if sufficient quantities of usable VOC's can
be recovered.
Flares are generally categorized in two ways: (1) by the
height of the flare tip (i.e., ground-level or elevated) and
(2) by the method of enhancing mixing at the flare tip (i.e.,
steam-assisted, air-assisted, pressure-assisted, or
unassisted). Elevating the flare can prevent potentially
dangerous conditions at ground level where the open flame
(i.e., an ignition source) is near a process unit. Further,
the products of combustion can be dispersed above working
areas to reduce the effects of noise, heat radiation, smoke,
and objectionable odors.
This discussion focuses on steam-assisted elevated
flares, the most common type used in the chemical industry.
Ground flares are discussed only in how they differ from
elevated flares.
The basic elements of a steam-assisted elevated flare
system are shown in Figure 2-1. The vent stream is sent to
the flare through the collection header (1). The vent stream
entering the header can vary widely in volumetric flow rate,
moisture content, VOC concentration, and heat content. The
knockout drum (2) removes water or condensed hydrocarbons that
can extinguish the flame or cause irregular combustion or
smoking in the flare combustion zone. Vent streams are also
typically routed through a flame arrester (3) before going to
the flare. This prevents possible flame flashback, which is
2-2
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Combustion Zone
(10)
Gas Collection Header
(1)
Vent Stream-*.
Knockout
Drum -*
(2)
Steam Nozzles
(9)
Flare Tip
(8)
Pilot Burners
(7)
Gas Barrier
(6)
Flare Stack
(5)
Purge
Gas 1
(4) A"
nr
V
Drain
Flame-*
Arrestor
(3)
Figure 2-1. Steam-assisted elevated flare system.
2-3
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caused when the vent stream flow rate to the flare is too low
and the flame front pulls down into the stack.1 Purge gas
(N2, CO2/ or natural gas) (4) also helps to prevent flashback
in the flare stack (5) caused by low vent stream flow rate.
To reduce the amount of purge gas required, a gas barrier (6)
or a stack seal is typically used just below the flare head to
impede the flow of air into the flare gas network.
The vent stream enters at the base of the flame where it
is heated by already burning fuel and pilot burners (7) at the
flare tip (8). Pilot burners positioned around the outer
perimeter of the flare tip ensure reliable ignition of the
vent stream. The flare tip is designed to give
environmentally acceptable combustion of the vent gas over the
flare system's capacity range. The maximum and minimum
capacity of a flare to burn a flared gas with a stable flame
(not necessarily smokeless) is a function of tip design. The
flare tip is designed to avoid a detached flame (a space
between the stack and flame with incomplete combustion), which
is caused by an excessively high flow rate.
The vent stream flows into the combustion zone (10),
where the exterior of the microscopic gas pockets are
oxidized. The rate of reaction is limited by the mixing of
the vent stream and oxygen from the air. If the gas pocket
has sufficient oxygen and residence time in the combustion
zone, complete combustion (all VOC's are converted to carbon
dioxide and water) will occur. Cracking can occur with the
formation of small hot particles of carbon that give the flame
its characteristic luminosity. If there is an oxygen
deficiency and if the carbon particles are cooled to below
their ignition temperature, smoking occurs.
A diffusion flame typical of elevated flares receives its
combustion oxygen by diffusion of air into the flame from the
surrounding atmosphere. The high flow rate of the vent stream
into the flare requires more combustion air at a faster rate
than simple gas diffusion can supply. Thus, many flares are
designed with "assisted" air supply and mixing. For example,
high-velocity steam injection nozzles (9) can be added to
2-4
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increase gas turbulence for mixing in the flame boundary
zones, thereby drawing in more combustion air and improving
combustion efficiency. This steam injection promotes
smokeless flare operation by minimizing the cracking reaction
that forms carbonaceous soot. The significant disadvantages
of using steam are increased noise and cost. The steam
requirement depends on the composition of the vent stream
flared, the velocity of the steam from the injection nozzle,
and the diameter of the flare tip. Although some gases can be
flared smokelessly without any steam, typically 0.01 to 0.6 kg
of steam per kilogram of vented gas is required.
Some flares use forced air instead of steam to provide
combustion air and the mixing needed for smokeless operation.
These flares consist of two coaxial flow channels. The
combustible gases flow in the center channel, and the
combustion air (provided by a fan in the bottom of the flare
stack) flows in the annulus. The principal advantage of
air-assisted flares is that they can be used where steam is
not available. Air assist is rarely used on large flares
because air flow is difficult to control when the gas flow is
intermittent. About 90.8 hp of blower capacity is required
for each 100 Ib/hr of gas flared.2
Ground flares are usually enclosed and have multiple
burner heads that are staged to operate based on the volume of
vent stream directed to the flare. The energy of the vent
stream itself (because of the high-pressure drop across the
nozzle) is usually adequate to provide the mixing necessary
for smokeless operation, and air or steam assist is not
needed. A fence or other enclosure reduces noise, heat, and
light from the flare and provides some wind protection.
Ground flares have less capacity than elevated flares and
are less widely used. Typically they are used to combust
continuous, constant flow vent streams, whereas steam-assisted
elevated flares are used to dispose of large amounts of gas
released in emergencies. Stable combustion can be obtained
with lower Btu-content vent streams than is possible with open
2-5
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flare designs (50 to 60 Btu/scf has been reported1) probably
due to their isolation from wind effects.
2.1.1.2 Factors Affecting Destruction Efficiency of
Flares. The destruction efficiency of flares is a function of
many factors: flammability, auto-ignition temperature, and
heat content of the vent stream; and mixing at the flare tip.1
The flammability limits of the vent stream flared
influence ignition stability and flame extinction.
Flammability limits are 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. That is, 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. It is easier to initiate and
maintain combustion of gases with wide limits of flammability
(for instance, H2).
The auto-ignition temperature of a vent stream affects
combustion because gas mixtures must be at a sufficient
temperature and concentration to burn. A gas with a low
auto-ignition temperature will ignite more easily than a gas
with a high auto-ignition temperature.
The heat content of the vent stream is a measure of the
heat available from the combustion of the VOC in the vent
stream. The heat content of the vent stream affects the flame
structure and stability. A gas with a lower heat content
produces a cooler flame that does not favor combustion
kinetics and is more easily extinguished. The lower flame
temperature will also reduce buoyant forces, which reduces
mixing.
Poor mixing at the flare tip or poor flare maintenance
can cause smoking (particulate matter release). Vent streams
with high carbon-to-hydrogen ratios (greater than 0.35) have a
greater tendency to smoke and require better mixing to burn
4
smokelessly. For this reason, one generic steam-to-vent-
stream ratio is not appropriate for all vent streams. The
2-6
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steam required depends on the vent stream carbon-to-hydrogen
ratio. A high ratio requires more steam to prevent a smoking
flare.
The efficiency of a flare in reducing VOC emissions can
be variable. For example, smoking flares are far less
efficient than properly operated and maintained flares.
Flares have been shown to have high VOC destruction
efficiencies, under proper operating conditions. Ninety-
eight percent combustion efficiency can be achieved by steam-
assisted flares with exit flow velocities less than 18 m/sec
(60 ft/sec) and combustion gases with heat contents over
11 MJ/scm (300 Btu/scf) and by flares operated without assist
with exit flow velocities less than 18 m/sec (60 ft/sec) and
burning gases with heat contents over 8 MJ/scm (200 Btu/scf).
2.1.1.3 Applicability of Flares. Flares can be
dedicated to a specific vent stream. Flares can also be
designed to control both normal process releases and emergency
upsets. The latter involves the release of large volumes of
gases. Often, large diameter flares designed to handle
emergency releases are also used to control continuous vent
streams from various process operations. Typically in
refineries, many vent streams are combined in a common gas
header to fuel boilers and process heaters. However, excess
gases, fluctuations in flow rate in the fuel gas line, and
emergency releases are sometimes sent to a flare.
Flares can be used to control almost any VOC stream and
can handle fluctuations in VOC concentration, flow rate, heat
content, and inerts content. Flaring is appropriate for
continuous, batch, and variable flow vent stream application.
Some streams, such as those containing halogenated or sulfur-
containing compounds, are usually not flared because they
corrode the flare tip or cause formation of secondary
pollutants (such as acid gases or sulfur dioxide).
2.1.2 Thermal Incinerators
2.1.2.1 Description of Thermal Incinerators. JLike other
combustion control devices, thermal incinerators operate on
the principle that any VOC heated to a high enough temperature
2-7
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in the presence of sufficient oxygen will be oxidized to
carbon dioxide and water. The theoretical combustion
temperature for thermal oxidation depends on the properties of
the VOC to be combusted. Some VOC's are oxidized at
temperatures much higher than others.
Thermal incineration processes are influenced by
residence time, mixing, and combustion temperature. An
efficient thermal incinerator system must provide:
• A chamber temperature high enough to enable
oxidation of the organic compounds to proceed
rapidly to completion;
• Sufficient turbulence for good mixing of the hot
combustion products from the burner, the combustion
air, and the organic compounds; and
• Sufficient residence time at the theoretical
combustion temperature for oxidation to reach
completion.
A thermal incinerator is usually a refractory-lined
chamber containing a burner (or set of burners) at one end.
As shown in Figure 2-2, discrete dual fuel burners (1) and
inlets for the vent stream (2) and combustion air (3) are
arranged in a premixing chamber (4) to mix the hot products
from the burners thoroughly with the process vent streams.
The mixture of hot combusting gases then passes into the main
combustion chamber (5). This chamber is sized to allow the
mixture enough time at the elevated temperature for oxidation
to reach completion (residence times of 0.3 to 1.0 sec are
common). Energy can be recovered from the hot flue gases in a
heat recovery section (6) to make the process more energy
efficient.6 Preheating combustion air or offgas is a common
method of energy recovery; however, it is sometimes more
economical to generate steam for use in the chemical process.
Insurance regulations require that if the vent stream is
preheated, the VOC concentration must be maintained below
25 percent of the lower explosive limit (LEL) to minimize the
potential for explosion hazards.
2-3
-------
Stack
Vent Stream
Inlet
(2)
Auxilllary
Burner
(Discrete)
Optional Heat
Recovery
(6)
Air
Inlet
(3)
Premixing
Chamber
(4)
V
Combustion
Chamber
(5)
Figure 2-2. Discrete burner, thermal incinerator.
2-9
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Thermal incinerators designed specifically for VOC
destruction using natural gas as the supplemental fuel may
also use a grid-type (distributed) gas burner (Figure 2-3).
The tiny gas flame jets (1) on the grid surface (2) ignite the
vapors as they pass through the grid. The grid acts as a
baffle to promote mixing of the gases entering the incinerator
chamber (3). This arrangement ensures that all vapors burn at
a lower chamber temperature, using less fuel. This system
makes possible a shorter reaction chamber, while maintaining
high efficiency.
Combustion devices are always operated with some quantity
of excess air to ensure a sufficient supply of oxygen. The
amount of excess air used varies with the fuel and burner type
but should be kept as low as possible to minimize fuel
consumption while maintaining combustion efficiency. Using
too much excess air wastes fuel because the additional air
must be heated to the combustion chamber temperature. Excess
air also increases flue gas volume and can increase the size
and cost of the system. Package, single-unit thermal
incinerators can be built to control streams with flow rates
in the range of 0.14 scm/sec (300 scfm) to about 24 scm/sec
(50,000 scfm). The smallest commercially available packaged
incinerators are designed to handle flow rates of 500 scfm.
Assuming a turndown ratio of 10 to I,8 the minimum flow rate
to an incinerator is 50 scfm.
2.1.2.2 Acid Gas Scrubbing. Thermal incinerators used
to reduce emissions of halogenated VOC's may require
additional control equipment to remove corrosive combustion
products (acid gas). The halogenated VOC streams are usually
scrubbed after combustion to prevent equipment corrosion that
results from contact with these acid gases. The flue gases
are quenched to lower their temperature and are then routed
through absorption equipment such as packed towers or liquid
jet scrubbers. Section 2.2.5 discusses the operation and
application of absorption equipment.
2.1.2.3 Thermal Incinerator Efficiency. Variations in
chamber temperature, residence time, inlet VOC concentration,
2-10
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Vent
Stream
inlet
Incinerator Chamber (3)
Burner Plate
Flame Jets
Auxiliary Fuel
(Natural Gas)
Stack
( Fan
Optional Heat
Recovery
(4)
Figure 2-3. Distributed burner, thermal incinerator,
2-11
-------
compound type, and flow regime (mixing) affect the VOC
destruction efficiency of a thermal incinerator. Performance
tests demonstrate that thermal incinerators can achieve
98 percent destruction efficiency for most VOC's at combustion
chamber temperatures ranging from 700 to 1,300 °C (1,300 to
2,370 °F) and residence times of 0.5 to 1.5 sec. For vent
streams with VOC concentrations below approximately
2,000 ppmv, all new thermal incinerators can achieve outlet
9
VOC concentrations of 20 ppmv or lower. These data indicate
that significant variations in destruction efficiency occurred
for GI to C§ alkanes and olefins, aromatics (benzene, toluene,
and xylene), oxygenated compounds (methyl ethyl ketone and
isopropanol), chlorinated organic compounds (vinyl chloride),
and nitrogen-containing compounds (acrylonitrile and
ethylamines) at chamber temperatures below 760 °C (1,400 °F).
A thermal incinerator properly designed and operated to
produce the described conditions in the combustion chamber
should be capable of higher than 98 percent destruction
efficiencies for any nonhalogenated VOC.
At temperatures above 760 °C (1,400 °F), oxidation occurs
more rapidly than gas diffusion mixing. The VOC destruction
efficiency then becomes dependent upon the fluid mechanics
(mixing) in the oxidation chamber. The flow regime must
ensure rapid, thorough mixing of the vent stream, combustion
air, and hot combustion products from the burner. This raises
the temperature of the VOC-laden stream and maintains it at
the combustion temperature, in the presence of excess oxygen,
for sufficient time to complete oxidation.
Other parameters affecting incinerator performance are
the heat content of the vent stream, the water content of the
stream, and the amount of excess combustion air (the amount of
air above the stoichiometric air needed for combustion).
Combustion of a vent stream with a heat content less than
1.9 MJ/scm (50 Btu/scf) usually requires burning supplemental
fuel to maintain the desired combustion temperature. Using
recuperative heat exchangers to preheat combustion air can
lessen or eliminate supplemental fuel requirements. Vent
i. .6
-------
streams with a heat content above 1.9 MJ/scm (50 Btu/scf) can
support combustion but may need supplemental fuel for flame
stability.
The maximum achievable VOC destruction efficiency
decreases with decreasing inlet VOC concentration because J
combustion is slower at lower inlet concentrations.
Therefore, a VOC weight percentage reduction based on the mass
rate of VOC exiting the control device versus the mass rate of
VOC entering the device is appropriate for vent streams with
VOC concentrations above approximately 2,000 ppmv
(corresponding to 1,000 ppmv VOC in the incinerator inlet
stream because air dilution is typically 1:1), however, for
lower inlet concentrations, an outlet concentration is more
appropriate than a percent reduction. As previously stated,
98 percent VOC reduction is achievable by incineration for
vent streams with VOC concentrations above 2000 ppmv, while an
incinerator outlet concentration of 20 ppmv is achievable for
vent streams with VOC concentrations below 2000 ppmv.
2.1.2.4 Applicability of Thermal Incinerators. ThermaJL_
incinerators, are technically feasible control devices for most
vent streams. They can be used for vent streams with any
concentration and type of VOC, and they can be designed to
handle minor fluctuations in flow rate. However, potential
excessive fluctuations in flow rate (i.e., process upsets)
might prevent the use of thermal incinerators and would
reguire the use of a flare. The presence of halogens reguires
additional eguipment such as scrubbers for acid gas removal.
2.1.3 Catalytic Incinerators
2.1.3.1 Description of Catalytic Incinerators. A
catalyst promotes oxidation of some VOC's at a lower
temperature than that required for thermal incineration. The
catalyst increases the rate of the chemical reaction without
becoming permanently altered itself. Catalysts typically used
for VOC incineration include platinum and palladium. Other
formulations used include metal oxides for emission streams
containing chlorinated compounds. Inert substrates are
2-13
-------
coated with thin layers of these materials to provide maximum
surface area for contact with the VOC in the vent stream.
Figure 2-4 is a schematic of a catalytic incineration
unit. The vent stream (1) is introduced into a mixing
chamber (2) where it is heated to about 316 °C (600 °F) by
contact with the hot combustion products from auxiliary
burners (3). The heated mixture is then passed through the
catalyst bed (4). Oxygen and VOC's migrate to the catalyst
surface by gas diffusion and are adsorbed to the surface of
the catalyst. Oxidation takes place at these active sites.
Reaction products are desorbed from the active sites and
transferred by diffusion back into the vent stream. The
combusted gas may then be passed through a waste heat recovery
device (5), such as a cross flow exchanger, to preheat the
incoming vent stream or combustion air before being discharged
to the atmosphere.
2.1.3.2 Catalytic Incinerator Control Efficiency.
Catalytic incinerators can achieve overall VOC desctruction
efficiencies up to about 98 percent and HAP desctruction
efficiencies up to about 95 percent with space velocities in
the range 30,000 to 100,000 gas hourly space velocity
(GHSV). The efficiency of the catalytic incinerator depends
on operating temperature, oxygen concentration, catalyst
activity, and the characteristics and concentration of the VOC
in the vent stream.
The operating temperatures of combustion catalysts
usually range from 316 to 650 °C (600 to 1,200 °F).
Temperatures below this range can slow or stop the oxidation
reaction. Higher temperatures can result in shortened
catalyst life and possible oxidation of the catalyst from the
support substrate.
The VOC content of the vent stream must be kept
relatively constant and low enough that the catalyst is not
overheated and its activity destroyed. To protect the
catalyst from overheating, VOC concentrations are usually
restricted to 25 to 30 percent of the LEL by insurance company
safety requirements. Such concentrations can be achieved by
2-14
-------
To Atmosphere
Stack
Auxiliary
Burners
(3)
Vent Stream
(1) -
Auxiliary
Burners
Catalyst Bed
Mixing Chamber
(2)
( Fan
Waste Heat
Recovery (optional)
Figure 2-4. Catalytic incinerator,
2-15
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diluting the vent stream with additional air if the
theoretical air to meet the stoichiometric oxygen requirements
for combustion is not sufficient.
Any accumulation of particulate matter, condensed VOC, or
polymerized hydrocarbons on the catalyst can block the
catalyst's effectiveness. Catalysts can also be deactivated
by compounds containing sulfur, bismuth, phosphorous, arsenic,
antimony, mercury, lead, zinc, tin, or halogens.13 If these
compounds exist in the vent stream, the VOC will pass through
the catalytic incinerator unreacted or be partially oxidized
to form compounds (aldehydes, ketones, and organic acids) that
are highly reactive atmospheric pollutants and can corrode
plant equipment.
2.1.3.3 Applicability of Catalytic Incinerators. The
applicability of catalytic incinerators for control of VOC's
is limited by the catalyst deactivation sensitivity to the
characteristics of the inlet stream. The vent stream to be
combusted should not contain materials that can poison the
catalyst or deposit on and block the reactive sites on the
catalyst surface. In addition, catalytic incinerators are
unable to handle high inlet concentrations of VOC or very high
flow rates. Catalytic incineration is generally useful for
concentrations of 50 to 10,000 ppmv, if the total
concentration is less than 25 percent of the LEL and for flow
rates of less than 100,000 scfm. Catalytic units are
typically used for vent streams with stable flow rates and
concentrations.
2.1.4. Industrial Boilers and Process Heaters
Industrial boilers and process heaters can be designed to
combust VOC's by incorporating the vent stream into the inlet
fuel or by feeding the stream into the boiler or heater
through a separate burner. The main distinction between
industrial boilers and process heaters is that the former
produces steam usually at high temperatures while the latter
raises the temperature of process streams.
2.1.4.1 Description of Industrial Boilers. JLndustriaJL
boilers are combustion units that boil ^ater to produce hjnh
,-16
-------
and low pressure steam. Industrial boilers can also be used
to combust various vent streams containing VOC's, including
vent streams from distillation operations, reactor processes,
and other general operations. This description focuses on the
use of industrial boilers to reduce emissions of VOC's from
vent streams.
The majority of industrial boilers used in the chemical
industry are of watertube design, and over half of these
boilers use natural gas as a fuel. In a watertube boiler,
hot combustion gases contact the outside of heat transfer
tubes which contain hot water and steam. These tubes are
interconnected by a set of drums that collect and store the
heated water and steam. The water tubes have relatively small
diameters, 5 cm (2.0 in.), which provide rapid heat transfer,
rapid response to steam demands, and relatively high thermal
efficiency. Energy transfer from the hot flue gases to the
water in the furnace watertube and drum system can be better
than 85 percent efficient. Additional energy can be recovered
from the flue gas by preheating combustion air in an air
preheater or by preheating incoming boiler feed water in an
economizer unit.
When firing natural gas, forced- or natural-draft burners
thoroughly mix the incoming fuel and combustion air. A VOC-
containing vent stream can be added to this mixture or it can
be fed into the boiler through a separate burner. In general,
burner design depends on the characteristics of the fuel—
either the combined VOC-containing vent stream and fuel or the
vent stream alone (when a separate burner is used).
2.1.4.2 Description of Process Heaters. A process
heater is similar to an industrial boiler in that heat
liberated by the combustion of fuels is transferred by
radiation and convection to fluids contained in tubular coils.
Process heaters are used in many chemical manufacturing
operations to drive endothermic reactions. They are also used
as feed preheaters and as reboilers for some distillation
operations. The fuels used in process heaters include natural
gas, refinery offgases, and various grades of fuel oil.
2-17
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Gaseous fuels account for about 90 percent of the energy
consumed by process heaters.17
There are many variations in the design of process
heaters depending on their application. In general, the
radiant section consists of the burner(s), the firebox, and
tubular coils containing the process fluid. Most heaters also
contain a convective section in which heat is recovered from
hot combustion gases by convective heat transfer to the
process fluid.
2.1.4.3 Factors Affecting Control Efficiency of
Industrial Boilers and Process Heaters. The combustion
efficiency of boilers and process heaters, like that of
incinerators, is determined by the average furnace temperature
and residence time. However, when a vent gas is injected as a
fuel into the flame zone of a boiler or process heater, the
required residence time is reduced because of the relatively
high temperature and turbulence of the flame zone.
Furnace residence time and temperature profiles vary for
industrial boilers and process heaters depending on the
furnace and burner configuration, fuel type, heat input, and
1W
excess air level. A mathematical model has been developed
that estimates the furnace residence time and temperature
19
profiles for a variety of industrial boilers. This model
predicts mean furnace residence times ranging from 0.25 to
0.83 sec for natural gas-fired watertube boilers that range in
size from 4.4 to 44 MW (15 to 150 MMBtu/hr). Boilers with a
44-MW capacity or greater generally have residence times and
operating temperatures that would ensure a 98 percent VOC
destruction efficiency. Furnace exit temperatures for this
size of boiler are at least 1,200 °C (2,200 °F), with peak
furnace temperatures in excess of 1,540 °C (2,810 °F).
Firebox temperatures for process heaters show relatively
wide variations depending on the application. In the chemical
industry, firebox temperatures can range from 400 °C (750 °F)
for preheaters and reboilers to 1,260 °C (2,300 °F) for
pyrolysis furnaces. Tests were conducted by the EPA to
determine trie benzene destruction efficiency ot ;.ive process
2-18
-------
heaters firing a benzene offgas and natural gas
mixture. ' ' The units tested are representative of
process heaters with low temperature fireboxes (reboilers) and
medium temperature fireboxes (superheaters). The reboiler and
superheater units tested showed greater than 98 percent
overall destruction efficiency for C^ to €5 hydrocarbons.
Additional tests conducted on a second superheater and a hot
oil heater showed that greater than 99 percent overall
destruction of C^ to Cg hydrocarbons occurred for both units.
2.1.4.4 Applicability of Industrial Boilers and Process
Heaters. Industrial boilers and process heaters are used
throughout the chemical industry to provide steam and heat
input essential to chemical processing. In most cases, these
industrial boilers and process heaters are of sufficient size
to provide the temperature and residence time needed for
destruction of VOC's. Further, industrial boilers and process
heaters have proved effective in destroying compounds that are
difficult to combust (e.g., polychlorinated biphenyls).
Therefore, industrial boilers and process heaters can be used
to reduce VOC emissions from any vent streams that are certain
not to reduce the performance or reliability of the boiler or
process heater.
The introduction of a vent stream into the boiler or
process heater flame zone can alter the heat transfer
characteristics of the furnace. Heat transfer characteristics
depend on the flow rate, heat content, and elemental
composition of the vent stream and on the size and type of
heat-generating unit used. Often, there is no significant
alteration of the heat transfer, and in some cases the VOC
content of the vent stream can reduce the amount of fuel
required to produce the desired heat release rate. In other
cases, the change in heat transfer characteristics after
introduction of a vent stream may affect the performance of
the heat-generating unit and increase fuel requirements. For
example, if the vent stream is a large-volume, low-VOC content
stream, additional energy may be needed to raise this stream
to the operating temperature of the boiler or process heater.
2-19
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Ducting some vent streams to a boiler or process heater
can present potential safety and operating problems. The
varying flow rate and organic content of some vent streams can
lead to explosive mixtures or flame instability within ^the
furnace.
Because they are corrosive, vent streams with halogenated
or sulfur-containing compounds are usually not combusted in
boilers or process heaters. When corrosive VOC's are
combusted, the flue gas temperature must be maintained above
the acid dew point to prevent acid deposition and subsequent
corrosion of downstream equipment such as ductwork, heat
exchangers, and exhaust stacks.
Boilers and process heaters are most applicable where the
potential exists for heat recovery from the combustion of the
vent stream. For example, vent stream? with a high VOC
concentration and high flow rate can provide enough equivalent
heat value to act as a substitute for fuel that would
otherwise be needed. Because boilers and process heaters
cannot tolerate wide fluctuations or interruptions in the fuel
supply, they are not widely used to reduce VOC emissions from
batch operations or other noncontinuous vent streams.
2.2 COLLECTION SYSTEMS AND RECOVERY DEVICES
This section discusses devices that reduce emissions of
VOC by means of product recovery. Because these devices
recover material without altering its chemical structure,
recovered feedstocks may be rpturned to the process and
recovered product may be collected and sold. A VOC is
recovered by separating it from the rest of the stream with a
process that is based on physical properties such as vapor
pressure or solubility.
The recovery devices discussed in the following five
subsections include loading rack vapor collection systems,
condensers, steam strippers, carbon adsorbers, and absorbers.
The discussion for each device treats its typical design
operation, factors affecting removal ef^icieno' »nri
applicabilitv.
-------
2.2.1 Vapor Collection Systems for Loading Racks
When liquids are transferred into a transport vessel,
vapors in the head space of that vessel can be lost to the
atmosphere. The principal factors affecting emissions from
transfer operations are the vapor pressure of the chemical
being transferred and the mode of transfer into the transport
vessel. Other factors that may influence emissions from
transfer operations include the transfer rate and the purge
rate of nitrogen (or other inert gas) through the vessel
during transfer.
The vapor pressure of the chemical being transferred has
the greatest influence on emissions from transfer operations.
For pure materials, the vapor pressure gives a measure of the
amount of organic compound lost during transfer. The total
potential emissions from any transfer is related to the void
volume of the transport vessel and the concentration of the
VOC in the head space. The saturation vapor pressure of the
pure VOC is used in calculating the maximum concentration
possible in the head space of the transport vessel. For low
volatility compounds, this vapor pressure is small and
therefore the emissions potential is small. For higher
volatility compounds, more of the compound evaporates into the
vapor space and is lost during transfer. For compounds with
vapor pressures above 14.7 psia, the liquid is usually
transferred under pressure, which effectively reduces the
potential for emissions.
The mode of transfer (loading) is also an important
factor in determining emissions from transfer operations. Top
splash loading creates the most emissions because it enhances
the agitation of the liquid being transferred, creating a
higher concentration of the compound in the vapor space. With
alternate loading techniques, such as submerged fill or bottom
loading, the organic liquid is loaded under the surface of the
liquid, which reduces the amount of agitation and suppresses
the generation of excess vapor in the head space of the
transport vessel.
2-21
-------
The transfer rate has a more subtle influence on
emissions; its greatest effect is on air quality. Transfer
rate will dictate the short-term emission rate of the compound
being transferred, thereby influencing exposure to the worker
or public.
A nitrogen purge is used to reduce the potential for
explosion of some chemicals in air or to keep some chemicals
moisture-free. Using an inert gas purge increases the
emission rate of VOC lost to the atmosphere because it creates
a turnover rate of gas through the transport vessel,
increasing the total volume of vapor discharged to the
atmosphere. If an inert gas purge is used, the total volume
of vapor loss is equal to the sum of the volume of the
transport vessel and the volume of the purge gas used. During
a transfer cycle, the volume of purge gas could be as great as
the volume of the transport vessel, essentially doubling the
losses to the atmosphere.
Most vapor collection systems collect the vapors
generated during transfer operations and transport them to
either a recovery device for return to the process or a
combustion device for destruction. In vapor balancing
systems, vapors generated during transfer operations are
returned directly to the storage facility for the material.
Assuming that incompressible fluids are being transferred,
negligible or no vapors would be lost to the atmosphere from
this "closed system," making no additional controls necessary.
2.2.1.1 Description of Vapor Collection Systems. Vapor
collection systems consist of piping that captures and
transports to a control device organic compounds in the vapor
space of a transport vessel that are displaced when liquids
are loaded. These systems may use existing piping normally
used to transport liquids under pressure into the transport
vessel or piping separate from that for transfer. Collection
systems comprise very few pieces of equipment and minimal
piping. The principal piece of equipment in a collection
system is a vacuum pump or blower, used to induce the flow of
2-22
-------
vapors from the transport vessel to the recovery or combustion
system.
Effective collection and recovery systems have been used
in some sectors of the SOCMI in response to Occupational
Safety and Health Administration (OSHA) regulations and EPA
standards. These systems transfer a VOC under pressure into
dedicated tank cars. Once the transfer has been completed,
the transfer piping is placed under vacuum with vapors being
drawn into a compression/refrigeration/condensation system.
The recovered liquid is stored for reuse, and the exhaust from
the recovery system is combusted in an on-site incinerator.
Once the transfer system has been completely evacuated, the
transfer piping is isolated by valves at the tank car and at
the transfer arm. The only potential emissions from this
system are the residual emissions from the combustion device
and the connection losses associated with the residual VOC in
the spool piece between the transfer arm and the transport
vessel. This system represents one of the most efficient
collection and recovery systems in the chemical industry.
Blowers can also be used to remove vapors from the head
space of the tank car as liquid is transferred into the tank
car. Standard recovery techniques such as condensation or
refrigeration/condensation systems, or combustion can be
applied to the captured vapors. The use of these lower
efficiency collection systems would depend upon the physical
properties of the chemicals transferred and on the efficiency
of the recovery device.
Vapor balancing is another means of collecting vapors and
reducing emissions from transfer operations. Vapor balancing
is most commonly used where storage facilities are adjacent to
the loading facility. As shown in Figure 2-5, an additional
line is connected from the transport vessel to the storage
tank to return any vapor in the transport vessel displaced by
the liquid that is loaded to the vapor space of the storage
vessel left by the transferred liquid. Because this is a
direct volumetric change, there should be no losses to the
atmosphere.
2-23
-------
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2-24
-------
2.2.1.2 Factors Affecting Efficiency of Vapor Collection
Systems. The factors affecting the efficiency of a vapor
collection system include the following:
• Operating pressure of the collection system;
• Volume of piping between the loading arm and the
transport vessel; and
• The efficiency of the ultimate control device.
The operating pressure of the collection system
influences the efficiency of collection through the
concentration of the VOC remaining in the line(s) after
transfer. For systems that operate at lower pressure or under
vacuum, the concentration of the VOC is decreased and the
total amount of VOC in the piping lowered. This effectively
reduces the amount of VOC that may be lost to the atmosphere
when disconnecting transfer lines. For systems operating at
higher pressures, there would be a larger quantity of VOC
remaining in the piping that could be lost.
The volume of piping between the transfer loading arm and
transport vessel, and the operating pressure of the collection
system establish the quantity of VOC not delivered to the
transport vessel and not collected for treatment. If the
piping is opened to the atmosphere, this quantity of VOC would
be lost, lowering the overall efficiency of the collection and
control system. Systems that minimize the piping between the
transfer loading arm and the transport vessel are more
"efficient" than those with larger piping connections.
The overall efficiency of the collection system and
control system is most dependent upon the efficiency of the
control system. In the SOCMI, collection systems are
generally hard-piped between the transport vessel and the
control system. Therefore, there would be no loss of
efficiency resulting from the collection system, other than
losses associated with connections and disconnections.
2-25
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2.2.1.3 Applicability of Vapor Collection Systems. The
applicability of and decision to use various vapor collection
systems depends upon several factors:
• Vapor pressure of the material;
• Value of the product;
• Physical layout of the facility; and
• OSHA considerations.
The vapor pressure of the material being transferred is a
principal consideration in the design of the vapor collection
system to be used. Materials with higher vapor pressures
(greater than 14.7 psia) are stored and loaded under pressure.
Loading under pressure eliminates the losses associated with
atmospheric transfer operations and limits losses to those
associated with connections and disconnections. Materials
with very low vapor pressures have relatively little emissions
potential and, therefore, little potential for emissions
reduction. Costs of collection and recovery systems for these
materials may be relatively high when considering the amount
and value of the recovered material. Systems such as vapor
balancing, however, would be an effective technique for
control of emissions from transfer operations of low vapor
pressure materials.
Proximity to the storage facility is the principal factor
in deciding to use vapor balancing. The viability of
installing a vapor balancing system is determined by the cost
of piping required to return the vapor to the storage tank.
When this distance is minimized, the cost is considered
affordable. Because vapor balancing is a simple and cost
effective control technique for transfer operations, it is
often viewed as a means of achieving reasonably available
control technology (RACT) requirements and has been used in
many instances as a control measure to meet the emission
requirements of State air toxic regulations. For purely
economic considerations, expensive products are candidates for
more extensive collection and recovery systems.
Additional considerations in the selection of a vapor
collection and recovery system are OSHA limitations on work
2-26
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place exposure to the chemicals being transferred. Some
chemical compounds have extremely strict threshold limit
values (TLV's) and, therefore, must be more tightly controlled
than other, less toxic compounds. The system described above
for the control of emissions from existing transfer operations
for some VOC's (using vacuum, vapor compression,
refrigeration, and combustion) is an example of extensive
control measures that limit exposure of individuals to the
chemical being transferred. Such systems have been used for
highly toxic or carcinogenic compounds.
2.2.2 Condensers
Condensation is a separation technique in which one or \
more volatile components of a vapor mixture are separated from
the remaining vapors through saturation followed by a phase /
change. The phase change from gas to liquid can be achieved
in two ways: (1) by increasing the system pressure at a given
temperature, or (2) by lowering the temperature at a constant
pressure. This section addresses the latter method.
In a two-component system where one of the components is
noncondensable (e.g., air), condensation occurs at dew point
(saturation) when the partial pressure of the volatile
compound is equal to its vapor pressure. For more volatile
compounds (i.e., compounds with lower normal boiling points),
a larger amount of the compound remains as vapor at a given
temperature; hence, to remove or recover the compound, a lower
temperature would be required for saturation and condensation.
For such cases, refrigeration can be used to obtain the lower
temperatures needed to achieve acceptable removal
efficiencies.
2.2.2.1 Description of Condensers. Figure 2-6 depicts a
typical configuration for a refrigerated surface condenser
system. The basic equipment includes a condenser,
refrigeration unit(s), and auxiliary equipment (e.g.,
precooler, recovery/storage tank, pump/blower, and piping).
The two most commonly used condenser types are surface
condensers and direct contact condensers. In surface
condensers, the coolant fluid does not contact the vent
2-27
-------
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2-28
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stream; heat transfer occurs through the tubes or plates in
the condenser. As the vapor condenses, a film forms on the
cooled surface and drains away to a collection tank for
storage, reuse, or disposal. Because the coolant from surface
condensers does not contact the vapor stream, it is not
contaminated and can be recycled in a closed loop. Surface
condensers also allow for direct recovery of VOC's from the
gas stream.
Most surface condensers in refrigerated systems are the
shell-and-tube type (Figure 2-7), which circulates the coolant
fluid on the tube side. The VOC's condense on the outside of
the tube (the shell side). Plate-type heat exchangers are
also used as surface condensers in refrigerated systems.
Plate condensers operate under the same principles as the
shell-and-tube systems (i.e., no contact between coolant and
vent stream), but the two streams are separated by thin, flat
plates instead of cylindrical tubes.
In contrast to surface condensers, direct contact
condensers cool the vapor stream by spraying a liquid at
ambient or lower temperature directly into the vent stream.
Spent coolant containing VOC's from direct contact condensers
usually cannot be reused directly. Additionally, VOC's in the
spent coolant cannot be recovered without further processing.
The combined VOC/coolant stream could present a potential
waste disposal problem, depending upon the coolant and the
specific VOC's.
For many VOC recovery needs, a refrigeration unit
generates the low-temperature medium necessary for heat
transfer. In refrigerated condenser systems, two kinds of
refrigerants are used—primary and secondary. Primary
refrigerants such as ammonia (R-717), and chlorofluorocarbons
such as chlorodifluoromethane (R-22) or dichlorodifluoro-
methane (R-12), are those that undergo a phase change from
liquid to gas after absorbing heat. Secondary refrigerants or
coolants, such as brine solutions, act only as heat carriers
and remain in the liquid phase. Conventional systems use a
closed primary refrigerant loop that cools the secondary
2-29
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(coolant) loop through heat transfer in the evaporator for the
primary refrigerant. The coolant is then pumped to a
condenser where it is used to cool the air/VOC vapor stream.
In some limited applications, however, the primary refrigerant
is used directly to cool the vapor stream.
For applications requiring low temperatures (below about
-30 °F), multistage refrigeration systems are frequently
employed. Multistage systems are of two types—compound and
cascade. In compound systems, refrigerant compression stages
are connected in series and 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 normally
considered for applications requiring temperatures between -50
to -150 °F and allow the use of different refrigerants in each
cycle. Theoretically, any number of cascaded stages are
possible, each stage requiring an additional condenser and an
additional stage of compression.
Some applications may require auxiliary equipment such as
precoolers, recovery/storage tanks, pumps/blowers, and piping.
If the vent stream contains water vapor 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 used to remove the moisture before
the vent stream enters the condenser. This precooler would
cool the vent stream to approximately 35 to 40 °F, effectively
removing the moisture from the vent stream. Alternatively,
ice can be melted during an intermittent heating cycle by
circulating ambient temperature brine through the condenser or
using radiant heating coils. If a system is not operated
continuously, ice can be removed by circulating ambient air.
A recovery tank for temporary storage of condensed VOC
before its reuse, reprocessing, or transfer to a large storage
tank may be necessary in some cases. Pumps and blowers are
2-31
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typically used to transfer liquid (e.g., coolant and recovered
VOC) and gas streams, respectively, within the system.
2.2.2.2 Factors Affecting Condenser Control Efficiency.
The removal efficiency of refrigerated surface condenser
systems designed to control vent streams containing air/VOC
mixtures depends primarily on the following parameters:
• Volumetric flow rate of the VOC-containing vent
stream;
• Inlet temperature of the vent stream;
• Concentrations of the VOC's in the vent stream;
• Absolute pressure of the vent stream;
• Moisture content of the vent stream; and
• Properties of the VOC's in the vent stream:
- dew points,
- heats of condensation,
heat capacities, and
vapor pressures.
Any component of any vapor mixture can be condensed if
brought to a low enough temperature and allowed to come to
equilibrium, but a condenser cannot lower the VOC
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 chlorofluorocarbons.
2.2.2.3 Applicability of Condensers. Condensers are
widely used as raw material and/or product recovery devices.
They may be used to recover VOC's upstream of other control
devices or they may be used alone for controlling vent streams
containing high VOC concentrations (usually greater than
5,000 ppmv). In these cases, the removal efficiencies of
condensers can range from 50 to 95 percent.
The temperature necessary for condensation depends on the
properties and concentration of VOC's in the vent stream.
Streams having low VOC concentrations and streams containing
more volatile (low boiling point) compounds require lower
condensation temperatures. Because condenser size and cost
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are a function of condensation temperature, condensation may
be economically infeasible for some vent streams.
Depending on the type of condenser used, disposal of the
spent coolant can be a problem. If cross-media impacts are a
concern, surface condensers would be preferable to direct
contact condensers.
Condensers used as emission control devices can process
flow rates as high as about 2,000 scfm. Condensers for vent
streams with flow rates above 2,000 scfm and high
concentrations of noncondensables will require significantly
larger heat transfer areas. These systems may be costly when
compared to the potential recovery of VOC or to other emission
reduction techniques.
2.2.3 Steam Strippers
Steam stripping can be used as initial treatment of a
process wastewater stream to reduce the VOC loading of that
stream before it is sent to the facility-wide wastewater
treatment system. A steam stripping system comprises several
components (including other devices described for control of
VOC's): a feed tank, heat exchanger, steam stripping column,
condenser, overheads receiver, and a destruction device (if
necessary).
2.2.3.1 Description of Steam Strippers. Steam stripping
involves the fractional distillation of wastewater to remove
VOC's. The basic operating principle of steam stripping is
the direct transfer of heat through contact of steam with
wastewater. This transfer of heat more easily vaporizes the
more volatile organic compounds. The overhead vapor that
contains water and organic compounds is condensed and
separated (usually in a decanter) to recover the organic
fraction. These recovered organic compounds are usually
either recycled for reuse in the process or incinerated in an
on-site combustion device for heat recovery.
Steam stripper systems may be operated in batch or
continuous mode. Batch steam strippers are more prevalent
when the wastewater feed is generated by batch processes, when
the characteristics of the feed are highly variable, or when
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small volumes of wastewater are generated. Batch strippers
may also be used if the wastewater contains relatively high
concentrations of solids, resins, or tars. With batch
stripping, wastewater to be steam stripped is charged to the
receiver, or pot, and brought to the boiling temperature of
the mixture. Heat is provided by direct injection of steam or
by an external heat exchanger, or reboiler. Solids, tars,
resins, and other residues remaining in the pot at the
completion of the batch are nonvolatile, heavy compounds that
are removed for disposal. Usually, batch steam strippers
provide a single equilibrium stage of separation. Therefore,
the removal efficiency is essentially determined by the
equilibrium coefficients of the pollutants and the fraction of
the initial charge distilled overhead. By varying the heat
input and fraction of the initial charge boiled overhead, a
batch stripper can be used to treat wastewater mixtures with
widely varying characteristics.
In contrast to batch strippers, continuous steam
strippers are designed to treat wastewater streams with
relatively consistent characteristics. Design of the
continuous stripper system is based on the flow rate and
composition of a specific wastewater feed stream or
combination of streams. Multistage, continuous strippers
normally achieve greater efficiencies of organic compound
removal than batch strippers. Continuous systems may offer
other advantages over batch stripping for applications
involving wastewater streams with relatively high flows and
consistent concentrations. These advantages include more
consistent effluent quality, more automated operation, and
lower annual operating costs.
Wastewater streams continuously discharged from process
equipment are usually relatively consistent in composition.
Such wastewater streams would be efficiently treated with a
continuous steam stripper system. However, batch wastewater
streams can also be controlled by continuous steam strippers
by incorporating a feed tank with adequate residence time to
provide a consistent outlet composition. Because the system
2-34
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can be designed to provide a stripper feed of consistent
quality, the remaining discussion focuses on continuous steam
stripping.
A generic continuous steam stripper system is shown in
Figure 2-8. For more effective control in specific cases,
alternate feed locations and multiple feed locations are
sometimes used. The steam stripper can also be operated under
a vacuum or may include a reflux stream where the bottoms
stream flows into a reboiler to vaporize, returning a portion
of the bottoms stream to the column. In addition, the pH of
the feed stream may be altered to change the equilibrium of a
low volatility compound, improving its removal through
stripping.
The purpose, design, and operation of each of the
components of a steam stripping system (a wastewater feed
tank, feed/bottoms heat exchanger and steam stripping column,
vent lines, condenser system, and ancillary pumps) are
discussed below in relation to their functions: collecting
and conditioning the wastewater, steam stripping of the
wastewater, controlling vents and openings in the system, and
recovery of the steam-stripped organic compounds.
2.2.3.1.1 Collecting and conditioning the wastewater.
The controlled sewer system or hard piping from the point of
wastewater generation to the feed tank controls emissions
before steam stripping. The feed tank, which is covered and
vented to an on-site combustion device, collects and
conditions the wastewater fed to the steam stripper. The feed
tank is usually sized to provide a desired hydraulic retention
time of 0.5 to 40 hr for the wastewater feed stream.26'27 The
desired retention time depends primarily on the variability in
wastewater flow rate, characteristics of the inlet wastewater,
and the amount of wastewater conditioning needed. Additional
surge capacity can provide retention time for wastewater
streams with highly variable flow rates to maintain a
relatively constant feed rate to the stripper.
If the feed tank is adequately designed, a continuous
steam stripper can treat wastewater generated by some batch
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2-36
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processes. In these cases, the feed tank serves as a buffer
between the batch process and the continuous steam stripper.
During periods of no wastewater flow from the batch process,
wastewater stored in the feed tank is fed to the stripper at a
relatively constant rate.
The retention time in the feed tank also depends on the
degree of conditioning required for the stripper feed stream.
Aqueous and organic phases are often present in the stripper
feed tank. The feed tank provides the retention time
necessary for these phases to separate. Oils and tars
normally partition from the water into the organic phase,
which is either recycled to the process for recovery of the
organic compounds or disposed by incineration. The water
phase is fed to the stripper to remove the soluble organic
compounds. Solids are also separated in the stripper feed
tank; the separation efficiency depends on the density of the
solids dissolved in the process wastewater. Some of the less
dense solids may remain suspended in the organic or aqueous
phases, while the more dense solids settle to the bottom of
the tank. These more dense solids are periodically removed
from the feed tank and are usually landfilled or landfarmed.
2.2.3.1.2 Steam stripping of the wastewater. After the
wastewater is collected and conditioned, it is pumped through
the feed/bottoms heat exchanger where it is preheated and then
pumped into the steam stripping column. Typically, steam is
sparged directly into the stripper at the bottom of the
column, and the wastewater feed is introduced into the
stripper at the top of the column (Figure 2-8). As the
wastewater flows down the column, it contacts the steam
flowing countercurrently up the column. Both latent and
sensible heat is transferred from the steam to the organic
compounds in the wastewater, vaporizing them into the gaseous
stream. These constituents flow out the top of the column
with any uncondensed steam.
The wastewater effluent leaving the bottom of the steam
stripper is pumped through the feed/bottoms heat exchanger
which heats the feed stream and cools the bottoms before
2-37
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discharge. This exchange of waste heat also improves the
economy of the steam stripper system. After passing through
the heat exchanger, the bottoms stream is usually either
routed to an on-site wastewater treatment plant (Figure 2-8)
and discharged to a National Pollutant Discharge Elimination
System (NPDES)-permitted outfall, or sent to a publicly owned
treatment works (POTW).
2.2.3.1.3 Controlling vents and openings in the steam
stripper system. In a steam stripper system, vent lines carry
gaseous organics, water vapor, and noncondensables to a
control device. For the stripper in Figure 2-8, vent lines
are placed between the stripper column and primary condenser,
between the primary condenser and feed tank, and between the
feed tank and existing on-site combustion device. All vent
lines are controlled with a combustion device (e.g., thermal
and catalytic incinerators, flares, boilers, or process
heaters) or a product recovery device (e.g., condensers,
carbon adsorbers, or absorbers).
Openings in a steam stripper system that allow for
pressure relief, venting, and/or maintenance access are sealed
unless in use. Pressure relief valves can be vented to a
control device. Because the typical stripper is operated at
pressures only slightly greater than atmospheric, pressure
relief emissions are typically negligible. Manways and other
access areas for maintenance and cleaning are never used
during operation and are sealed to prevent emissions during
operation.
2.2.3.1.4 Recovery of the steam stripped organic
compounds. A condenser system can be used to recover the
organic and water vapors in the gaseous overheads stream from
the stripping column. If the primary condenser is not
sufficient for condensing a large portion of the organic
compounds, it may be necessary to install a secondary
condenser with brine or a refrigerant. The condensed
overheads stream is fed to an overhead receiver. The
recovered organic compounds are usually either pumped to
storage and recycled to the process unit or combusted for
2-38
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their fuel value in an incinerator, boiler, or process heater.
These combustion devices are discussed in Section 2.1. If an
aqueous phase is generated in the overhead receiver, it is
returned to the feed tank and recycled through the steam
stripper system. Any noncondensable gases and highly volatile
organic compounds not recovered by the condenser system are
routed to an on-site control device such as a carbon adsorber,
boiler, process heater, or an incinerator. Often in practical
systems, the vent from the overhead condenser is piped to the
feed tank, which is vented to the control device.
2.2.3.2 Factors Affecting Steam Stripper Control
Efficiency. Steam stripping achieves emission reductions of
0 to 99 percent, based on the chemical characteristics (e.g.,
strippability) of the wastewater stream. However, 95 to
99 percent reduction can be achieved for the majority of
compounds regulated by the HON. The organic compound removal
performance of the steam stripper depends on the degree of
contact between the steam and the wastewater. Several factors
affecting the degree of contact that occurs in the steam
stripper column are: (1) the dimensions of the column (height
and diameter); (2) the contacting media in the column (trays
or packing); and (3) operating parameters such as the steam-
to-feed ratio, column temperatures, and pH of the wastewater.
Increasing the column height increases removal efficiency
by increasing the opportunity for contact between the steam
and the wastewater. The column height is determined by the
number of theoretical stages required to achieve the desired
removal efficiency. The number of theoretical stages is a
function of the equilibrium coefficient of the pollutants and
the efficiency of mass transfer in the column.
The diameter of the column is determined from the
required cross-sectional area for liquid and vapor flow
through the column. High superficial gas velocities increase
turbulence and mixing and result in high column efficiencies.
However, the column cross-sectional area must be sufficient to
prevent flooding from excessive liquid loading or liquid
entrainment due to excessive vapor velocity. The cross-
2-39
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sectional area also affects the liquid retention time in the
column, with higher retention times resulting in higher
efficiencies. These opposing factors are considered in
selecting the design values of superficial gas and liquid
velocities (flow per square unit of tower cross-sectional
area) and the tower diameter.
The contacting media in the column play a major role in
determining the mass transfer efficiency. Typically, steam
stripping columns are equipped with trays (e.g., bubble-cap or
sieve) or packing to provide contact between the vapor and
liquid phases. Trays are regularly spaced throughout the
column and provide staged contact between the two phases;
packing provides for continuous contact. Trays are usually
more effective for wastewater containing dispersed solids
because of the plugging and cleaning problems encountered with
packing. Also, tray towers can operate efficiently over a
wider range of liquid flow rates than packed towers. Packed
towers are often more cost effective to install and operate
when treating highly corrosive wastewater streams because
corrosion resistant ceramic packing can be used. Also, the
pressure drop through packed towers may be less than the
pressure drop through tray towers designed for equivalent
wastewater loadings. Packed towers are seldom designed with
diameters greater than 4 ft, and column heights may be more
limited than that of tray towers due to crushing of the
packing near the bottom of the column. However, with the use
of intermediate packing supports, crushing problems can be
reduced.
An increase in the steam-to-feed ratio will increase the
ratio of the vapor to liquid flow through the column. This
increases the stripping of organic compounds into the vapor
phase. Because additional heat is provided when the steam
rate is increased, additional water is also vaporized.
Therefore, an increase in the steam-to-feed ratio is also
accompanied by an increase in the steam rate flowing out of
the column in the overheads stream. Steam-to-feed ratios
generally range from 0.01 to 1.0 kg/kg.
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The steam-to-feed ratio required for high removal
efficiency is affected by the temperature of the wastewater
fed to the column. If the feed temperature is lower than the
operating temperature at the top of the column, part of the
steam is required to heat the feed. Good column design
provides sufficient steam flow to heat the feed and volatilize
the organic constituents. Steam in excess of this flow rate
helps to carry the organic compounds out of the top of the
column with the overheads stream.
Column operating temperature and wastewater pH also
affect organic compound removal performance. Temperature
affects the solubility and equilibrium coefficients of the
organic compounds. Stripping columns are usually operated at
pressures slightly greater than atmospheric, and operating
temperatures are usually slightly greater than the normal
boiling point of water. Wastewater pH is often controlled by
adding caustic to the feed to raise the pH and change the
vapor-liquid equilibrium characteristics of the compound in
the wastewater matrix and enhance the removal of organic
po
compounds being steam stripped.
Removal efficiency of a steam stripper decreases with low
inlet organic concentrations. Above an inlet concentration
threshold of 50 to 100 ppm, the removal efficiency for most
compounds is not affected by the concentration of organic
compounds in the wastewater; however, at concentrations well
below this threshold, steam strippers lose some effectiveness.
Finally, organic compound removal efficiencies are higher
for steam strippers treating wastewater containing chlorinated
organic compounds. These chlorinated organic compounds are
more easily steam stripped than organic compounds such as
phenol which are more soluble in water and less volatile.
2.2.3.3 Applicability of Steam Strippers. Based on
industry responses to Section 114 information requests,
10 percent of the reported wastewater streams account for
greater than 90 percent of the volatile organic HAP compounds
by mass. In addition, 20 percent of the total reported
wastewater flow accounts for 85 percent of the volatile
2-41
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29
organic HAP compounds by mass. Therefore, it may be
possible to achieve significant VOC emission reduction by
controlling a relatively small number of individual wastewater
streams and a relatively small volume of wastewater. In many
cases, it may be possible to combine two or more of these
streams for treatment by the same steam stripper. By hard
piping these streams from the point of generation to the steam
stripper, emissions from transporting the wastewater are
minimized. As streams are combined, the cost of control
increases, but the cost of control per stream decreases. This
issue is discussed further in the presentation of steam
stripper costs in Chapter 3.
Steam stripping is most applicable to treating
wastewaters with organic compounds that are highly volatile
and have a low solubility in water. The VOC's that have low
volatility tend to volatilize less readily and thus are not
easily stripped out of the wastewater by the steam.
Similarly, VOC's that are very soluble in water tend to remain
in the wastewater and also are not easily stripped out by
steam. There is also an interactive effect between these two
properties; for example, methanol is fairly volatile but has a
high solubility in water. Steam stripping wastewaters
containing methanol removes approximately one third of the
actual methanol content. At least one steam stripper studied
operated at an altered pH to change the vapor-liquid
equilibrium characteristics of a less volatile compound. The
lower-volatility compound was removed efficiently through this
pH change.
Oil, grease, and solids content and pH of a wastewater
stream also affect the applicability of steam stripping to a
particular wastewater. High levels of oil, grease, and solids
can cause operating problems for steam strippers. High or low
pH may prove to be corrosive to equipment. However, these
problems can usually be circumvented by design or wastewater
preconditioning techniques.
2-42
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2.2.4 Carbon Adsorbers
2.2.4.1 Description of Carbon Adsorbers. Adsorption is \
a mass-transfer operation involving interaction between
gaseous- or liquid-phase components and solid-phase
/
components. This section addresses the adsorption of VOC's
from vent streams. The gas phase (adsorbate) is captured on
the solid-phase (adsorbent) surface by physical or chemical
adsorption mechanisms. Physical adsorption takes place when
intermolecular (van der Waals) forces attract and hold the gas
molecules to the solid surface. Chemisorption occurs when a
chemical bond forms between the gaseous- and solid-phase
molecules. A physically adsorbed molecule can be removed
readily from the adsorbent (under suitable temperature and
pressure conditions); the removal of a chemisorbed component
is much more difficult.
The most common industrial adsorption systems use
activated carbon as the adsorbent. Activated carbon
effectively captures certain organic vapors by physical
adsorption. The vapors can then be released for recovery by
regenerating the adsorption bed with steam or nitrogen.
Compared to activated carbon, oxygenated adsorbents such as
silica gels, diatomaceous earth, alumina, or synthetic
zeolites exhibit a greater selectivity for capturing water
vapor than organic gases. Thus, these oxygenated adsorbents
would be of little use for the high-moisture vent streams
characteristic of some VOC-containing vent streams.
The design of a carbon adsorption system depends on the
chemical characteristics of the VOC being recovered, the
physical properties of the inlet stream (temperature,
pressure, and volumetric flow rate), and the physical
properties of the adsorbent. The mass of VOC that adheres to
the adsorbent surface is directly proportional to the
difference in VOC concentration between the gas phase and the
solid surface. In addition, the quantity of VOC adsorbed
depends on the adsorbent bed volume, the surface area of
adsorbent available to capture VOC, and the rate of diffusion
of VOC through the gas film at the gas- and solid-phase
2-43
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interface (the mass transfer coefficient). Physical
adsorption is an exothermic operation that is most efficient
within a narrow range of temperature and pressure.
Figure 2-9 is a schematic diagram of a typical fixed-
bed, regenerative carbon adsorption system. Before it enters
the carbon bed, the inlet vent stream is usually filtered to
prevent bed contamination by soot, resin droplets, and large
particles entrained in the vent stream and cooled to maintain
the bed at optimum operating temperature and to prevent fires
or polymerization of the hydrocarbons (1). The filtered and
cooled vapors enter the adsorber stage of the system (2),
where they pass through the porous-activated carbon bed.
The dynamics of the process may be illustrated by viewing
the carbon bed as a series of layers or mass-transfer zones
(3a, b, c). Gases entering the bed are highly adsorbed first
in Zone a. Because most of the VOC is adsorbed in Zone a,
very little adsorption takes place in Zones b and c.
Adsorption in Zone b increases as Zone a reaches equilibrium
with organic compounds. The process continues until Zone c is
used. When the bed is completely saturated (breakthrough),
the incoming VOC-laden stream is routed to an alternate bed
while the saturated carbon bed is regenerated.
The carbon bed is regenerated by heating the bed or
applying vacuum to draw off the adsorbed gases. Low pressure
steam (4) is frequently used as a heat source to strip the
adsorbent (carbon) of the VOC collected. After steaming, the
carbon bed is cooled and dried, typically by blowing air
through it with a fan, and the steam-laden vapors are routed
to a condenser (5) and on to a VOC recovery system (6). The
regenerated bed is returned to active service while the
saturated bed is purged of VOC. The regeneration process may
be repeated numerous times; eventually, however, the carbon
must be replaced.
2.2.4.2 Factors Affecting Carbon Adsorption Control
Efficiency. For well designed and operated carbon adsorber
systems, continuous VOC removal efficiencies of more than
95 percent are achievable for a variety of solvents, including
2-44
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VOC-Laden
Vent Stream
(1)
Fan
Low
Pressure (4)
Steam
Closed
Open ,
-M2
Adsorber 1
(adsorbing)
Steam
Adsorber 2
(regenerating)
Open
Closed
(5)
Condenser
Vent to
Atmosphere
Decanter and/or
Distilling Tower
(6)
Recovered
"*" Solvent
Water
Figure 2-9. Two-stage regenerative adsorption system.
2-45
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mixtures that contain ketones such as methyl ethyl ketone and
cyclohexanone. Several plants have been shown to continuously
achieve removal efficiencies of 97 to 99 percent.31 The VOC
removal efficiency of an adsorption unit depends on the inlet
vent stream characteristics, the physical properties of the
compounds present in the vent stream, the physical properties
of the adsorbent, and the condition of the regenerated carbon
bed.
Inlet stream temperature, pressure, and velocity affect
adsorption unit efficiency. The adsorption capacity of the
carbon and the resulting outlet concentration are dependent
upon the temperature of the inlet vent stream. High vent
stream temperature increases the kinetic energy of the gas
molecules, causing them to overcome van der Waals forces and
release from the surface of the carbon. At vent stream
temperatures above 100 °F, both adsorption capacity and outlet
concentration may be adversely affected. In these cases,
inlet stream coolers are typically used.
Increasing vent stream pressure improves VOC removal
efficiency. Increased stream pressure results in higher
concentration of VOC in the vapor phase and an increased
driving force for mass transfer to the carbon surface.
Conversely, carbon beds are often regenerated by inducing low
pressure or a vacuum. Reducing the pressure in the carbon bed
effectively lowers the concentration of VOC's in the vapor
phase, desorbing the VOC's from the carbon surface to the
vapor phase.
The velocity of the vent stream entering the carbon bed
must be quite low to allow time for diffusion and adsorption.
Linear velocities of 50 to 100 fpm are typically used. At
higher velocities the pressure drop across the carbon bed
becomes too high for standard blowers. At lower velocities,
the bed becomes too large and expensive. If inlet
concentrations are low (as is expected in the SOCMI) the bed
area required for the volume of carbon needed usually permits
a velocity at the high end of this range.
2-46
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The required depth of the bed for a given compound is
directly proportional to the carbon granule size and porosity
and to the inlet vent stream velocity. For a given carbon
type, bed depth must increase as the vent stream velocity
increases. Generally, carbon adsorber bed depths range from
1.5 to 3.0 ft. A bed depth of at least 1.5 ft is used to
ensure that the bed is substantially deeper than the mass
. _ 34
transfer zone.
The condition of the carbon bed will change with use.
After repeated regeneration, the carbon bed loses activity,
resulting in reduced VOC removal efficiency.
2.2.4.3 Applicability of Carbon Adsorption. Although
carbon adsorption is an excellent method for recovering some
valuable process chemicals, it cannot be used universally for
distillation or process vent streams. Carbon adsorption is
not recommended under the following conditions, which are
common with many VOC-containing vent streams: (1) high VOC
concentrations, (2) very high or low molecular weight
compounds, (3) mixtures of high and low boiling point VOC's,
and (4) high moisture content.
The maximum practical inlet concentration of VOC is
typically 10,000 ppmv. Adsorbing vent streams with high VOC
concentration may result in excessive temperature rise in the
carbon bed due to the accumulated heat of adsorption resulting
from the VOC loading. If flammable vapors are present,
insurance company requirements may limit inlet concentrations
to less than 25 percent of the LEL. However, vent streams
with high VOC concentrations can be diluted with air or inert
gases to make a workable adsorption system.
The molecular weight of the compounds to be adsorbed
should be in the range of 45 to 130 gm/gm-mole for effective
adsorption. High molecular weight compounds that are
characterized by low volatility are strongly adsorbed on
carbon. The affinity of carbon for these compounds makes it
difficult to remove them during regeneration of the carbon
bed. Hence, carbon adsorption is generally not applied to
such compounds (i.e., boiling point above 400 °F and molecular
2-47
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weight greater than about 130). Because highly volatile
materials (i.e., molecular weight less than about 45) do not
adsorb readily on carbon, adsorption is not typically used for
controlling emission streams containing such compounds.
Properly operated adsorption systems can be very
effective with homogeneous vent streams but less so with a
stream containing a mixture of light and heavy hydrocarbons.
The lighter organic compounds tend to be displaced by the
heavier (higher boiling) components, greatly reducing system
efficiency.
At adsorbate concentrations above 1,000 ppmv, humidity
does not significantly affect working capacity. However, at
adsorbate concentrations below 1,000 ppmv, water vapor
competes with the VOC's in the vent stream for adsorption
sites on the carbon surface. In these cases, vent stream
humidity levels exceeding 50 percent (relative humidity) are
not desirable. Moisture content may be reduced by adding
drier dilution air to the vent stream or by passing the stream
38
through a heat exchanger.
2.2.5 Absorbers
2.2.5.1. Description of Absorbers. Absorption is the
selective transfer of one or more components of a gas mixture
into a liquid solvent. The transfer consists of solute
diffusion and dissolution into a solvent. For any given
solvent, solute, and operating conditions, there exists an
equilibrium ratio of solute concentration in the gas mixture
to solute concentration in the solvent. The driving force for
mass transfer at a given point in an operating absorber is the
difference between the concentration of solute in the gas and
the equilibrium concentration of solute in the liquid. If the
solvent contains an additive that will react with the solute,
absorption will be enhanced because the equilibrium
concentration of solute in the liquid will be zero.
Devices based on absorption principles include spray
towers, venturi and wet impingement scrubbers, packed columns,
and plate columns. Spray towers require high atomization
pressure to obtain solvent droplets 500 to 100 urn in size,
2-48
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necessary for a sufficiently large surface area for contact
between the liquid and gas streams.39 Although they can
remove particulate matter effectively, spray towers have the
least effective mass transfer capability and are generally
restricted to particulate matter removal and control of
high-solubility gases such as sulfur dioxide and ammonia.40
Venturi scrubbers have a high degree of gas/liquid mixing and
provide high particulate matter removal efficiency. They also
require high pressure drops (i.e., they are energy intensive)
and have relatively short contact times. Therefore, their use
41
is also restricted to high-solubility gases. As a result,
VOC control by gas absorption is generally accomplished in
packed or plate columns.
Packed towers are vertical columns containing inert
packing, manufactured from materials such as porcelain, metal,
or plastic, that provides the surface area for contact between
the liquid and gas phases in the absorber. Packed columns are
used primarily for corrosive materials and liquids with
tendencies to foam or plug. They are less expensive than
plate columns for small-scale or pilot plant operations where
the column diameter is less than 0.6 m (2 ft). They are also
suitable where the use of plate columns would result in
excessive pressure drops.
In plate or tray columns, contact between the gas and
liquid phases takes place in a stepwise fashion on a series of
trays. The liquid phase flows down from tray to tray as the
gas phase moves up through openings (e.g., perforations or
bubble caps) in the tray, passing through the liquid along the
way. These columns are preferred for large-scale operations,
where internal cooling is desired or where low liquid flow
rates would inadequately wet the packing.42
Figure 2-10 is a schematic of a packed tower using
countercurrent flow. The vent stream containing VOC's to be
absorbed is introduced near the bottom of the tower (1) and
allowed to rise through the packing material (2). Solvent
flows in from the top of the column, countercurrent to the
vapors (3), absorbing the solute from the gas phase and
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toF
Absorbing
Liquid In
(5) Cleaned Gas Out
inal Control Device
or to Atmosphere
; Packing (2)
(1) VOC-Laden
Gas In
Absorbing, Liquid
with VOC Out
to Disposal or VOC/Solvent Recovery
Figure 2-10. Packed tower absorption process.
2-50
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carrying the dissolved solute out of the tower (4). The vent
stream after treatment exits at the top (5) for release to the
atmosphere or for further treatment as necessary. The
saturated liquid may be sent to a stripping unit where the
absorbed VOC is recovered. After the stripping operation the
absorbing solution is either recycled to the absorber or sent
to a water treatment facility for disposal.
The major tower design parameters to be determined for
absorbing any substance are column diameter and height, system
pressure drop, and liquid flow rate required. These
parameters are derived by considering the solubility,
viscosity, density, and concentration of the VOC in the inlet
vent stream (all of which depend on column temperature); the
total surface area provided by the tower packing material; and
the quantity of gases to be treated.
2.2.5.2. Factors Affecting Absorption Control
Efficiency. Control efficiencies for absorbers vary from 50
to 95 percent. The VOC removal efficiency of an absorber
depends on the solvent selected and on proper design and
operation. Absorbing liquids (solvents) are chosen for high
solubility for the specific VOC and include liquids such as
water, mineral oils, kerosenes, nonvolatile hydrocarbon oils,
and aqueous solutions of oxidizing agents, sodium carbonate,
43
and sodium hydroxide. For a given solvent and solute, an
increase in absorber size (i.e., contact surface area) or a
decrease in the operating temperature can increase the VOC
removal efficiency of the system. It may be possible in some
cases to increase VOC removal efficiency by a change in the
solvent.
2.2.5.3. Applicability of Absorption. The applicability
of absorption for controlling VOC emissions is dependent
primarily upon the availability of a suitable solvent. Water
may be used for the absorption of organic compounds that have
relatively high water solubilities (e.g., most alcohols,
organic acids, aldehydes, ketones, amines, and glycols). For
organic compounds that have low water solubilities, other
2-51
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solvents (usually organic liquids with low vapor pressures)
, 44
are used.
Other factors bearing on the applicability of absorption
as a viable emissions control option include the absorptive
capacity and strippability of the VOC in the solvent. These
factors are especially important to regenerable systems.
Absorptive capacity is a measure of the solubility of the VOC
in the solvent. The solubility limits the total quantity of
VOC that could be absorbed in the system. Strippability
relates to the ease with which the VOC can be removed from the
solvent. If a VOC cannot be easily desorbed from the solvent,
then absorption is a less viable control technique for control
of the VOC.45
The VOC concentration in the inlet vent stream also
determines the applicability of absorption. Absorption is
usually considered only when the VOC concentration is above
44
200 to 300 ppmv. Below these gas-phase concentrations, the
rate of mass transfer of the VOC to the solvent is decreased
sufficiently to make reasonable designs infeasible.
2.3 STORAGE TANK IMPROVEMENTS FOR EMISSION REDUCTION
This section discusses two types of vessels commonly used
for storage in the chemical manufacturing industry: fixed
roof tanks and internal floating roof tanks.
Both types of tanks are cylindrical and have an axis
perpendicular to the ground. Each tank has a cone- or dome-
shaped roof that is permanently affixed to the tank shell.
The internal floating roof tank also has a roof inside the
tank that floats on the surface of the stored liquid. A
conservation vent, which is a type of pressure- and vacuum-
relief valve, is commonly installed on fixed roof tanks to
control losses from the tank caused by minor changes of
pressure in the vapor space.
The major types of emissions from fixed roof tanks are
breathing and working losses. Breathing loss is the expulsion
of vapor from a tank vapor space because of diurnal changes in
temperature and barometric pressure. Emissions occur in the
absence of any change in the liquid level in the tank.
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Working losses occur during filling when the liquid level
in the tank rises. The vapors are expelled from the tank when
the pressure inside the tank exceeds the relief pressure as a
result of filling.
Loss of VOC from internal floating roof tanks occurs in
one of four ways:
(1) Through the annular rim space around the perimeter
of the floating roof (rim or seal losses);
(2) Through the openings in the deck required for
various types of fittings (fitting losses);
(3) Through the nonwelded seams formed when joining
sections of the deck material (deck seam losses);
and
(4) Through evaporation of liquid left on the tank wall
following withdrawal of liquid from the tank
(withdrawal loss).
As wind flows over the exterior of an internal floating
roof tank, air flows into the enclosed space between the fixed
and floating roofs through some of the shell vents and out of
the enclosed space through others. Any VOC that has
evaporated from the exposed liquid surface and that has not
been contained by the floating deck will be swept out of the
enclosed space.
Emissions from storage tanks can be reduced by reducing
the surface area of liquid that is exposed to air. This can
be achieved by: installing a floating roof inside a fixed
roof tank, or upgrading an existing internal floating roof.
The following sections discuss these tank improvements as
emission control techniques. Their potential for emission
reduction is compared to fixed roof tanks, which are
considered the minimum acceptable equipment currently used for
the storage of volatile organic liquids (VOL's).
2.3.1 Description of Tank Improvements
The presence of a floating roof (deck) inside a fixed
roof tank significantly reduces the surface area of exposed
liquid. The deck serves as a physical barrier between the VOL
and the air that enters the tank through vents. _^
2-53
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To retrofit a fixed roof tank with an internal floating
deck, the tank must be cleaned and degassed and vents must be
cut in the tank roof to minimize the possibility of
hydrocarbon vapors accumulating in flammable concentrations.
Installing sections of a welded deck in the tank will require
cutting an opening in the tank shell larger than any standard
manway.
Upgrading an existing floating roof involves cleaning and
degassing the tank and replacing some parts of the deck (e.g.,
rim seals and deck fittings) with parts that allow fewer
emissions. For example, uncontrolled deck fittings
(Table 2-1) are replaced with controlled fittings. Liquid-
mounted rim seals may be substituted for vapor-mounted seals.
Also, a secondary rim seal may be added to supplement the
control achieved by the primary seal. Access to the tank for
upgrading an existing internal floating roof is usually made
through an existing manway.
2.3.2 Factors Affecting Control Efficiency
Because evaporation is the primary emission mechanism
associated with storage tanks, emissions from both fixed roof
and floating roof tanks vary with the vapor pressure of the
stored liquid. Thus, the control efficiency of retrofitting a
fixed roof tank with an internal floating deck depends upon
the material being stored.
Other factors affecting emissions, and therefore control
efficiency, are tank size, number of turnovers, and the type
of deck and seal system selected. Installing an internal
floating roof can reduce emissions by 76 to 98 percent.4 The
relative effectiveness of one internal floating roof design
over another is a function of how well the deck can be sealed.
To simplify the discussion, an example tank is used as a
common basis for evaluating effectiveness. The example tank
has the following characteristics:
• Tank diameter—85 ft;
• Tank height—47 ft;
• Tank capacity—2,000,000 gal;
• Vapor pressure of the stored liquid—0.71 psia;
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TABLE 2-1.
CONTROLLED AND UNCONTROLLED INTERNAL
FLOATING ROOF DECK FITTINGS
Equipment Description
Deck Fitting Type
Uncontrolled
Controlled
1. Access hatch
2. Automatic gauge
float well
3. Column well
4. Ladder well
5. Sample pipe or
well
6. Vacuum breaker
Unbolted, ungasketed
cover3; or unbolted,
gasketed cover
Unbolted, ungasketed
cover3; or unbolted,
gasketed cover
Built-up column-
sliding cover,
ungasketed3; pipe
column-sliding
cover, ungasketed
Ungasketed sliding
cover3
Sample well with
slit fabric seal,
10% open area3
Weighted mechanical
actuation,
ungasketed3
Bolted, gasketed
cover
Bolted, gasketed
cover
Flexible fabric
sleeve seal or a
gasketed sliding
cover
Gasketed sliding
cover
Sample well with
slit fabric
seal, 10% open
area3
Weighted
mechanical
actuation,
gasketed
3Fittings assumed in the uncontrolled case for estimating the
effectiveness of fittings controls. This scenario is
representative of no single tank, but rather is the composite
of what is estimated based on a survey of users and
manufacturers to be typical of fittings on the majority of
tanks currently in service. Note that the sample well with
split fabric seal was used in the uncontrolled case for
calculating emissions because it is in common use. It was
also used in the controlled case because it is the lowest
emitting fitting type.
2-55
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• Density of stored liquid—7.22 Ib/gas;
• Molecular weight of the product vapor—92.13 Ib/lb-
mole; and
• Turnover rate--seven per year.
Probably the most typical internal floating roof design
is the noncontact, bolted, aluminum internal floating roof
with a single vapor-mounted wiper seal and uncontrolled
fittings. As discussed above, there are four types of losses
from this roof design. Contributions made by the four types
of losses to the total loss from the example tank are
estimated as follows:
(1) Rim or seal losses: 33 percent;
(2) Fitting losses: 35 percent;
(3) Deck seam losses: 30 percent; and
(4) Withdrawal losses: 2 percent.
With the exception of withdrawal losses, which are
inherent in all internal floating roof designs, the losses
listed above can be reduced by employing roofs with
alternative design features. Table 2-2 lists alternative
floating roof equipment designs and the emission rate
associated with each type of equipment installed on the
example tank. The following sections, organized according to
types of losses, elaborate on the alternative equipment that
can be employed on internal floating roofs.
2.3.2.1 Control Of Seal losses. Internal floating roof
seal losses can be minimized by employing liquid-mounted
primary seals instead of vapor-mounted seals and/or by
employing secondary wiper seals in addition to primary seals.
Available emissions test data suggest that the location
of the seal (i.e., vapor- or liquid-mounted) and the presence
of a secondary seal are the major factors affecting seal
losses. A liquid-mounted primary seal has a lower emissions
rate, and thus a higher control efficiency, than a vapor-
mounted seal. Equipping the example tank with an internal
floating roof having a liquid-mounted rim seal results in an
emission reduction of 96.0 percent over a fixed roof tank;
whereas a deck with a vapor-mounted seal achieves only
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TABLE 2-2. EFFECTIVENESS OF INTERNAL FLOATING ROOFS
ON AN EXAMPLE TANKa
Internal Floating Roof Tank
Example Fixed
Roof Tank
Total emission
rate
17.78 Mg/yr
(working loss
= 9.90)
(breathing
loss = 7.88)
Case
1
2
3
4
5
6
7
8
Equipment
Typeb/c
fclFRvm
bIFRvm,cf
bIFRvm,cf ,ss
bIFRlm,ss
bIFRlm,cf
bIFRlm,cf ,ss
wIFRlm,cf ,ss
Total
Emission
Rated
(Mg/yr)
0.88
0.76
0.58
0.71
0.65
0.60
0.54
0.29
Reduction
Over
Fixed Roof
Tank
Emission
Rate (%)
95.1
95.7
96.7
96.0
96.3
96.6
97.0
98.4
aExample tank is 2,000,000-gal capacity; 85 ft in diameter,
47 ft in height, 0.71 psia vapor pressure, 92.13 Ib/lb-mole
molecular weight of product and condensed product vapor, and
seven turnovers per year.
bKey:
b = bolted roof deck; w = welded roof deck; IFR = internal
floating roof; v = vapor-mounted primary seal; m = metallic
shoe primary seal; 1 = liquid-mounted primary seal;
cf = controlled fittings; ss = rim-mounted secondary seal.
Uncontrolled fittings (as designated by absence of cf):
(1) access hatch, with ungasketed, unbolted cover;
(2) automatic gauge float well, with ungasketed, unbolted
cover; (3) built-up column well, with ungasketed sliding
cover; (4) ladder well, with ungasketed sliding cover;
(5) adjustable roof legs; (6) sample well with slit fabric
seal (10 percent open area); (7) 1-in. diameter stub
drains; and (8) vacuum breaker with gasketed weighted
mechanical actuation.
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TABLE 2-2. EFFECTIVENESS OF INTERNAL FLOATING ROOFS
ON AN EXAMPLE TANK3 (CONCLUDED)
Controlled fittings (cf): (1) access hatch, with gasketed,
bolted cover; (2) automatic gauge float well, with gasketed,
bolted cover; (3) built-up column well, with gasketed,
sliding cover; (4) ladder well, with gasketed sliding cover;
(5) adjustable roof legs; (6) sample well with slit fabric
seal (10 percent open area); (7) 1-in. diameter stub
drains; and (8) vacuum breaker, with gasketed weighted
mechanical actuation.
dReference 47.
47
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95.1 percent reduction (Table 2-2). A secondary seal, with
either a liquid- or a vapor-mounted primary seal, provides an
additional level of control.
The type of seal used plays a less significant role in
/ «
determining the emissions rate. The type of seal is
important only to the extent that the seal must be suitable
for the particular application. For instance, an elastomeric
wiper seal is commonly employed as a vapor-mounted primary
seal or as a secondary seal for an internal floating roof.
Because of its shape, this seal is not suitable for use as a
liquid-mounted primary seal. Resilient filled seals, on the
other hand, can be used as both liquid- and vapor-mounted
seals.
The effectiveness of alternate internal floating roof
seal systems can be evaluated by inspecting the rim seal loss
factors (Kr). These factors have been developed based on test
data for estimating losses for various seal systems and are
listed in Table 2-3 for seals with average gaps. Estimates
of control efficiency and incremental control efficiency are
also listed in the table. The control efficiency estimates
indicate the effectiveness of the various seal systems at
reducing emissions over the level achieved by a vapor-mounted
primary seal. In Table 2-3, the vapor-mounted primary seal is
assumed to be the baseline control level. The incremental
control efficiency estimates demonstrate the effectiveness of
each seal system relative to the next less stringent seal
system (i.e., the next higher emitting seal system). These
efficiencies are calculated directly from the Kr values.
As shown in the table, application of a liquid-mounted
primary and secondary seal system in place of a vapor-mounted
primary seal would reduce seal losses an estimated 76 percent.
On the example tank, where these seal losses represent roughly
one-third of the total loss from the tank (i.e., bIFRvm case),
this 76 percent reduction in seal losses translates to a
25 percent reduction in the total loss from the floating roof
tank. Relative to fixed roof tank emissions, the additional
control provided by the liquid-mounted primary and secondary
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TABLE 2-3. INTERNAL FLOATING ROOF RIM SEAL SYSTEMS,
SEAL LOSS FACTORS, AND CONTROL EFFICIENCIES
Seal System (Ib-mole/ft-yr)
Seal Loss
Control
Efficiency
Related
To Baseline
Incremental
Control
Efficiency
Vapor-mounted
primary seal
only
Liquid-mounted
primary seal
only
Vapor-mounted
primary and
secondary
seals
Liquid-mounted
primary and
secondary
seals
6.7
3.0
2.5
1.6
IFRb baseline
(0%)
55%
63%
55%
17%
76%
36%
aKr = Rim seal loss factors.
bIFR = Internal floating roof type.
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seal system over the vapor-mounted primary seal system
increases the effectiveness of the internal floating roof from
95.1 percent to 96.3 percent for the example tank. (See
Case 1 vs. Case 5 in Table 2-2.)
2.3.2.2 Control Of Fitting Losses. The numerous
fittings that penetrate or are attached to an internal
floating roof include access hatches, column wells, roof legs,
sample pipes, ladder wells, vacuum breakers, and automatic
gauge float wells. Fitting losses occur when VOC's leak
around these fittings. Fitting losses can be controlled with
gasketing and sealing techniques or by the substitution of
fittings that are designed to leak less. Table 2-1 lists the
fitting types considered representative of uncontrolled and
controlled fittings.
The effectiveness of fitting controls at reducing the
overall emission rate is a function of the number of fittings
of each type employed on a given tank. On the example tank,
use of controlled fittings reduces the total fitting loss by
about 36 percent. Because fitting losses are about 35 percent
of the total emissions from the example internal floating roof
tank (i.e., for an IFR^), the controlled fittings reduce the
overall emissions by about 13 percent over a similar tank
without fitting controls. The additional emission reduction
achieved by installing controlled fittings increases the
control efficiency of the example internal floating roof tank
from 95.1 percent to 95.7 percent over a fixed roof tank as
the base case.
2.3.2.3 Control Of Deck Seam Losses. Depending on the
type of floating roof used, deck seam losses can contribute to
the total loss from an internal floating roof. For the
example tank used as a basis for comparison throughout this
section (i.e., blFRvm)/ deck seam losses are 30 percent of the
total loss. When seal losses and fitting losses are
controlled, the relative contribution to the total loss from
deck seams increases. In the case of a bolted, noncontact,
internal floating roof with liquid-mounted primary seals,
controlled deck fittings, and secondary seals (bIFRlm,cf,ss)/
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deck seam losses contribute about 47 percent of the total
loss.
Deck seam losses are inherent in several floating roof
types. Any roof constructed of sheets or panels fastened by
mechanical fasteners (e.g., bolts) is expected to have deck
seam losses. Two roof types were tested to determine deck
49
seam losses. The first was a bolted, aluminum, noncontact
roof and the second was a bolted, aluminum panel-type, contact
roof. The design of the mechanical fasteners used on these
two roof types varies significantly. In addition, one roof
type floats above the liquid surface while the other floats in
contact with the liquid surface. Despite these differences,
the seams on these two roof types were found to emit at
roughly the same rate per meter of seam. Deck seam losses,
therefore, are considered to be a function of the length of
the seams and not the type of mechanical fastener or the
position of the deck relative to the liquid surface.
Deck seam losses are controlled by selecting a roof type
with vapor-tight deck seams. The welded deck seams on steel
pan roofs are vapor tight. Fiberglass-lapped seams of a glass
fiber reinforced polyester roof may be vapor tight as long as
there is negligible permeability of the liquid through the
seam-lapping materials. Some manufacturers provide gaskets
for bolted metal deck seams. One vendor of contact, aluminum
bolted decks has provided bench-scale test results that
substantiate emissions reduction due to gasketing deck
50
seams.
Selecting a welded roof (rather than a bolted roof) will
eliminate deck seam losses. On an internal floating roof with
liquid-mounted primary seal, secondary seal, and controlled
fittings, the elimination of deck seam losses improves the
overall effectiveness for the example tank from a 97.0 to
98.4 percent control efficiency (see Case 7 vs. Case 8 in
Table 2-2), relative to a fixed roof tank.
2.3.3 Applicability of Storage Tank Improvements
The applicability of storage tank improvements as a
technology to reduce VOC emissions is dependent upon
2-62
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characteristics of the particular VOC. Because floating decks
are often constructed of aluminum panels, this control
technology may not be applicable to tanks storing halogenated
compounds, pesticides, or other compounds that are
incompatible with aluminum. Contact between these compounds
and an aluminum deck could corrode the deck and cause product
contamination.
In addition, chemical vapor pressure may affect the
selection of tank improvements as an applicable control
technology. For chemicals with very low vapor pressure, fixed
roof tank emissions will already be so low that installing an
internal floating roof may not significantly reduce emissions
further. (This lower vapor pressure limit cannot be specified
because the equations for calculating emissions from internal
floating roof tanks are only applicable for vapor pressures
above 0.1 psia). For chemicals with medium vapor pressures
5"v
(up to about 9 psia), emission reductions of 95 percent and
above are achievable with this technology. However, for
chemicals with higher vapor pressures, achievable emission
reduction starts to decrease with increasing vapor pressure.
Thus, tank improvements may not have acceptable emission
reduction efficiency for chemicals with high vapor pressures.
2.4 EQUIPMENT LEAK EMISSION SOURCES AND EMISSION CONTROL
TECHNIQUES
Equipment leak emissions in the chemical manufacturing
industries occur when process fluid (either liquid or gaseous)
is released through the sealing mechanisms of equipment in the
plant. This section discusses the sources of equipment leak /
emissions and control techniques that can be applied to reduce
emissions from equipment leaks, including the applicability of
each control technique and its associated effectiveness in
reducing emissions.
There are many potential sources of equipment leak
emissions in an organic chemical processing unit. The
following sources of equipment leaks are considered: pumps,
compressors, agitators, pressure relief devices, open-ended
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lines, sampling connections, process valves, connectors,
instrumentation systems, and product accumulator vessels.
Emissions from equipment leaks have been the subject of
intense study by the EPA and industry. Emissions from the
populations of equipment leak sources have been found to be
random occurrences, when considering individual pieces of
equipment. Further, no relationship between emissions and
process variables, such as line temperatures and pressure, has
been established. Precise quantitation of emission reduction
effectiveness for the techniques described in this section has
not been made on a unit-by-unit basis, but approximate
quantitation has been made on an industry-wide basis for
certain emission reduction techniques. Many of these were
described in detail in the EPA's additional information
document on emissions, emissions reductions, and costs of
controlling emissions from equipment leaks.51
The techniques for reducing emissions from equipment
/ leaks are as diverse as the types of sources. These
\ techniques can be classified in the following major
\categories:
\ • Equipment (modifications and leakless equipment);
• Closed vent systems; and
• Work practices.
The selection of a control technique and its effectiveness in
reducing emissions depends on a number of factors, including:
• Type of equipment (valve, pump, compressor, etc.);
• Equipment service (gas, light liquid, heavy liquid);
• Process variables that limit equipment selection
(temperature, pressure);
• Process stream composition; and
• Costs.
The various control techniques that can be applied to the
different types of equipment components and their estimated
potential emission reductions are presented in the following
sections. Section 2.4.1 describes the various equipment
components that can be sources of emissions and the equipment
modifications and alternative equipment that can be applied to
2-64
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reduce emissions from these equipment components. Closed vent
systems and work practices are described in Sections 2.4.2 and
2.4.3, respectively. Both closed vent systems and certain
work practices can be applied to a number of different
equipment types and therefore are described in more general
terms.
2.4.1 Equipment Description and Specifications
Equipment leaks may be reduced by retrofitting with
equipment designed to reduce or prevent leakage. Equipment
modifications and leakless equipment for each equipment type
are described below. Some of the modifications may be
applicable to more than one type of equipment.
2.4.1.1 Pumps. Pumps are used extensively in the
chemical industries for the movement of organic liquids.
The centrifugal pump is the most widely used pump in the
chemical industries; however, other types, such as the
positive displacement (reciprocating) pump, are also used.
Chemicals transferred by pump can leak at the point of contact
between the moving shaft and the stationary casing.
Consequently, all pumps except the sealless type, such as
canned-motor, magnetic drive, and diaphragm pumps, require a
seal at the point where the shaft penetrates the housing in
order to isolate the pumped fluid from the environment.
2.4.1.1.1 Seals for pumps. Two generic types of seals,
packed and mechanical, are used on pumps. Packed seals can be
used on both reciprocating and rotary action (centrifugal)
pumps. A packed seal consists of a cavity ("stuffing box") in
the pump casing filled with packing material that is
compressed with a packing gland to form a seal around the
shaft. Coolant is required to remove the frictional heat
between the packing and shaft. The necessary lubrication is
provided by a coolant that flows between the packing and the
shaft. Deterioration of the packing can result in leakage of
process liquid.
Mechanical seals are limited in application to pumps with
rotating shafts. There are single and double mechanical
seals, with many variations to their basic design, but all
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have a lapped seal face between a stationary element and a
rotating seal ring.53 The rotating seal ring and stationary
element faces are lapped to a very high degree of flatness to
maintain contact throughout the entire mutual surface area.
In a single mechanical seal, the faces are held together by
the pressure applied by a spring on the drive and by the pump
pressure transmitted through the pumped fluid on the pump end.
An elastomer O-ring seals the rotating face to the shaft. The
stationary face is sealed to the stuffing box with another
elastomer O-ring or gasket.
For double mechanical seals, two seals are arranged back-
to-back, in tandem, or face to face. In the back-to-back
arrangement, a closed cavity is created between the two seals.
A seal liquid, such as water or seal oil, is circulated
through the cavity. Because the seal liquid surrounds the two
seals, it can be used to control the temperature in the
stuffing box. For the seal to function properly, the pressure
of the seal liquid must be greater than the operating pressure
of the pump. In this manner, any leakage would occur across
the seal faces into the process or the environment.
Therefore, the seal liquid must be compatible with the process
54
fluid and the environment.
In a tandem dual mechanical seal arrangement, the seals
face the same direction. The secondary seal backs up the
primary seal. The cavity between the two seals is filled with
a barrier fluid, which may serve to control temperature of the
seals. In the tandem arrangement, the barrier fluid may be
maintained at a pressure lower than that in the stuffing box,
allowing any leakage through the seal from the process into
the barrier fluid. To reduce emissions from such an
.arrangement, a barrier fluid reservoir system should be used.
At the reservoir, the process liquid can vaporize (i.e.,
degas) and be released to the atmosphere, or to a control
device for proper recovery or destruction.
In the face-to-face arrangement, two rotating faces are
mated with a common stationary element. The barrier fluid in
the cavity between the seals may be maintained at higher or
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lower pressures than the stuffing box. As in the tandem
arrangement, leakage from the process fluid to the barrier
fluid can occur if the barrier fluid is maintained at a
pressure lower than that in the stuffing box. Preventing
emissions from the barrier fluid requires (1) operation of the
barrier fluid at a pressure higher than that of the stuffing
box or (2) use of a degassing reservoirs vented to a control
device.
The actual emission reduction achievable through use of
double mechanical seals depends on the frequency of seal
failure. Failure of both the inner and outer seals can result
in relatively large emissions at the seal area of the pump.
Pressure monitoring of the barrier fluid may be used to detect
failure of the seals. Visual inspection of the seal area
can also detect failure of the outer seals. Seal failure
would require the leaking pump to be shut down for
maintenance.
Double mechanical seals are used in many process
applications; however, there are some conditions that preclude
their use. Their maximum service temperature is usually less
than 260 °C, and mechanical seals can rarely be used
successfully on pumps with reciprocating shaft motion.
Further, double mechanical seals cannot be used where the
process fluid contains slurries, polymeric, or undissolved
solids.
2.4.1.1.2 Sealless pumps. Another type of pump used in
the chemical industry is the sealless pump (i.e., canned-
motor, diaphragm, and magnetic drive pumps). Sealless pumps
are used primarily in processes where the pumped fluid is
hazardous, highly toxic, or very expensive and where every
effort must be made to prevent all possible leakage of the
fluid. Canned-motor pumps have interconnected cavity
housings, motor rotors, and pump casings. As a result, the
motor bearings run in the process liquid and all seals are
eliminated. Because the process liquid is the bearing
lubricant, abrasive solids in the process lines cannot be
tolerated. Canned-motor pumps are widely used for handling
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organic solvents, organic heat transfer liquids, and light
oils.
Diaphragm pumps perform similarly to piston and plunger
pumps. However, the driving member is a flexible diaphragm of
metal, rubber, or plastic. The primary advantage of this
arrangement is the elimination of all packing and seals
exposed to the process liquid provided the diaphragm's
integrity is maintained. This is important when handling
hazardous or toxic liquids. Emissions from diaphragm pumps
can be large, however, if the diaphragm fails.
In magnetic-drive pumps, no seals contact the process
fluid. An externally-mounted magnet coupled to the pump motor
drives the impeller in the pump casing.
2.4.1.2 Compressors. Compressors provide motive force
for transporting gases through a process unit in much the same
way that pumps transport liquids. Compressors are typically
driven with rotating or reciprocating shafts. Thus, the
sealing mechanisms for compressors are similar to those for
pumps, that is, packed and mechanical seals. Emissions from
this source type may be reduced by improving the seals'
performance or by collecting and controlling the emissions
from the seal. Emissions from mechanical contact seals depend
on the type of seal or control device used and the frequency
of seal failure.
Shaft seals for compressors are of several different
types: labyrinth, restrictive carbon rings, mechanical
contact, and liquid film. All of these seal types restrict
leaks, although none of them completely eliminates leakage.
Compressors can be equipped with ports in the seal area to
evacuate collected gases, which could then be controlled.
The labyrinth compressor seal is composed of a series of
close-tolerance, interlocking "teeth," which restrict the flow
of gas along the shaft. Labyrinth seals have the largest leak
potential of the different seal types. Many variations in
"tooth" design and materials of construction are available to
improve compressor seal performance. Properly applied "tooth"
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configuration and shape can reduce leakage by up to 40 percent
over a straight pass labyrinth seal.56
Restrictive carbon ring seals consist of multiple
stationary carbon rings with close shaft clearances. This
type of seal may be operated dry, with a sealing fluid, or
with a buffer gas. Restrictive ring seals can achieve lower
leak rates than labyrinth seals.
Mechanical contact seals are similar to the mechanical
seals described for pumps. Clearance between the rotating and
stationary elements is reduced to essentially zero. Oil or
another suitable lubricant is supplied to the seal faces.
Mechanical contact seals can achieve the lowest emission rates
of the types of compressor seals described here. Like
mechanical seals for pumps, however, they are not suitable for
all processing conditions.
A buffer or barrier fluid may be used with these
mechanical seals to form a buffer between the compressed gas
and the environment, similar to barrier fluids in pumps. This
system requires a clean, external gas supply that is
compatible with the gas being compressed. Barrier gas can
become contaminated and must be disposed of properly, for
example by venting to a control device.
Centrifugal compressors also can be equipped with liquid
film seals. The seal is formed by a film of oil between the
rotating shaft and stationary gland. Process gas can be
released to the atmosphere when the circulating oil is
returned to the oil reservoir. To eliminate release of
volatile air pollutants from the seal oil system, the
reservoir can be vented to a control device.
2.4.1.3 Agitators. Agitators are used in the chemical
industries to stir or blend chemicals. As with pumps and
compressors, emissions from agitators can occur at the
interface of a moving shaft and a stationary casing.
Emissions from this source type may be reduced by improving
the seal or by collecting and controlling emissions. Four
seal arrangements are commonly used with agitators: packed
seals, mechanical seals, hydraulic seals, and lip seals.
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Packed seals for agitators are very similar in design and
application to the packed seals for pumps (Section 2.4.1.1).
Although mechanical seals are more costly than other seal
arrangements, they provide better leakage rate reduction.
Also, the maintenance frequency of properly installed and
maintained mechanical seals is one-half to one-fourth that of
packed seals. In fact, at pressures greater than 1,140 kPa
(165 psia), packed seals are rarely used because the
performance of mechanical seals is superior. 9 Mechanical
seals can be designed specifically for these higher pressure
applications. As with packed seals, the mechanical seals for
agitators are similar to the design and application of
mechanical seals for pumps.
The hydraulic seal is the simplest and least-used
agitator shaft seal. In this type of seal, an annular cup
attached to the process vessel contains a liquid that contacts
an inverted cup attached to the rotating agitator shaft. The
primary advantage of this seal is that it is a noncontact
seal. However, this seal is limited to low temperatures and
pressures and can only handle very small pressure
fluctuations. Process chemicals may contaminate the seal
liquid and then be released into the atmosphere as equipment
leak emissions.
Lip seals, which are relatively inexpensive and easy to
install, can be used on a top-entering agitator as a dust or
vapor seal. The sealing element in the lip seal is a spring-
loaded elastomer. Once the seal has been installed, the
agitator shaft rotates in continuous contact with the lip
seal. Pressure limitations of the seal are 14 to 21 kPa (2 to
3 psi) because it operates without lubrication. Operating
temperatures are limited by the characteristics of the
selected elastomer. Emissions can be released through this
seal when it wears excessively or when the operating pressure
surpasses the pressure limitation of the seal.
2.4.1.4 Pressure Relief Devices. Insurance, safety, and
engineering codes require that pressure relief devices or
systems be used in applications where the process pressure may
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exceed the maximum allowable working pressure of the process
equipment. Pressure relief devices include rupture disks and
safety/relief valves. The most common pressure relief device
is a spring-loaded valve designed to open when the operating
pressure of a piece of process equipment exceeds a set
pressure. In this example, the pressure relief valve is
constructed so that it will close or reseat after the
operating pressure has decreased to a level below the set
pressure. Equipment leak emissions from spring-loaded relief
valves may be caused by failure of the valve seat or valve
stem, improper reseating after overpressure relief, or process
operation near the relief valve set pressure which may cause
the relief valve to frequently open and close or "simmer."
Manufacturers of relief valves state that resilient seat
or "O-ring" relief valves provide better reseat qualities
after an overpressure relief compared to standard relief
valves. The applicability of resilient seat technology is
limited by material compatibility and operating conditions.
This technology would have no effect on emissions from
overpressure episodes, emissions due to failure of the seat
surfaces, or "simmering."
Rupture disks are designed to burst at overpressure to
allow the process gas to vent directly to the atmosphere.
Rupture disks allow no emissions as long as the integrity of
the disk is maintained. The rupture disk must be replaced
after each pressure relief episode to restore the process to
an operating pressure condition. Although rupture disks can
be used alone, they are sometimes installed upstream of a
relief valve to prevent emissions through the relief valve
seat.
Rupture disk/relief valve combinations require certain
design constraints and criteria to avoid potential safety
hazards. For example, appropriate piping changes must be made
to prevent disk fragments from lodging in and damaging the
relief valve when relieving overpressure. A block valve
upstream of the rupture disk can be used to isolate the
rupture disk/relief valve combination and permit in-service
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replacement of the disk after it bursts. Otherwise, emissions
could result through the relief valve.
A dual-pressure relief system with a three-way valve can
be used to maintain operation while having continuous
overpressure protection. The three-way valve can be open to
either one of the dual rupture disk (RD)/pressure relief
valves (PRVs) so that one RD/PRV combination is always in
service. If one of the rupture disks fails, it can be made
available for repair by switching the three-way valve to the
other RD/PRV combination. Rupture disk/relief valve
combinations must have some provision for testing the
integrity of the disk. The American Petroleum Institute (API)
has provided guidance on the design of instrumentation to
indicate failure of the rupture disk. According to API
guidelines, the area between the rupture disk and relief valve
should be connected to a pressure indicator, recorder, or
alarm. The control efficiency of the disk/valve combination
is assumed to be 100 percent when operated and maintained
properly. If disk integrity is not maintained or if the disk
is not replaced after overpressure relief, the control
efficiency is lowered, but this reduction in efficiency has
not been quantified.
2.4.1.5 Open-Ended Lines. Emissions from open-ended
lines are caused by leakage through the seat of an upstream
valve in the open-ended line. Emissions that occur through
the stem and gland of the valve are not considered "open-
ended" emissions and are addressed in Section 2.4.1.7.
Emissions from open-ended lines can be controlled by
installing a cap, plug, flange, or second valve to the open
end. The control efficiency of these control measures is
assumed to be 100 percent. In the case of a second valve, the
upstream valve should always be closed first after use of the
valves to prevent the trapping of fluids between the valves.
Each time the cap, plug, flange, or second valve is opened,
any VOC that has leaked through the first valve seal will be
released.
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2.4.1.6 Sampling Connections. Emissions from sampling
connections occur as a result of purging the sampling line to
obtain a representative sample of the process fluid. Based on
survey data on the SOCMI, approximately 25 percent of the
valves on open-ended lines are used for sampling
connections. Emissions from sampling connections can be
reduced by using a closed loop sampling system or disposing of
the purged process fluid in a control device. The closed loop
sampling system is designed to return the purged fluid to the
process at a point of lower pressure. A throttle valve or
other device is used to induce the pressure drop across the
sample loop. Closed loop sampling is assumed to be
100 percent effective for controlling emissions from sample
purge. The purged fluid could also be directed to a control
device such as an incinerator, in which case the control
efficiency would depend on the efficiency of the incinerator
in removing the VOC. Because some pressure drop is required
to purge sample through the loop, low pressure processes or
storage tanks may not be amenable to closed loop sampling.
Also, safety requirements may prohibit closed loop sampling in
some instances.
2.4.1.7 Process Valves. Valves are the most common and
numerous process equipment type found in the chemical
industry. Valves are available in many designs, most of which
contain a valve stem that operates to restrict or allow fluid
flow. Typically, the stem is sealed by a packing gland or
O-ring to prevent leakage of process fluid to the atmosphere.
Emissions from valves occur at the stem or gland area of the
valve body when the packing or O-ring in the valve fails.
2.4.1.7.1 Seals for valves. Valves that require the
stem to move in and out or turn must utilize a packing gland.
A variety of packing materials are suitable for conventional
packing glands. The most common packing materials are the
various types of braided asbestos that contain lubricants;
other packing materials include graphite, graphite-impregnated
fibers, and tetrafluorethylene. The choice of packing
material depends on the valve application and configuration.62
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Conventional packing glands can be used over a wide range of
operating temperatures. At high pressures, these glands must
be quite tight to attain a good seal.63
Elastomeric 0-rings are also used for sealing process
valves, but are not suitable where there is sliding motion
through the packing gland. These seals are rarely used in
high-pressure service, and operating temperatures are limited
by the seal material.64
2.4.1.7.2 Sealless valves. Emissions from process
valves can be eliminated if the valve stem can be isolated
from the process fluid. Two types of leakless or sealless
valves are available: diaphragm valves and sealed bellows
valves.
Diaphragm valves isolate the valve stem from the process
fluid using a flexible elastomer or metal diaphragm. The
position of the diaphragm is regulated by a plunger, which is
controlled by the stem. The stem may be actuated manually or
automatically by standard hydraulic, pneumatic, or electric
actuators. In this arrangement, the stem/plunger pushes the
diaphragm toward the valve bottom, throttling the flow of the
process fluid. The diaphragm and stem/plunger are connected
so that it is impossible for them to be separated under normal
working conditions. When the diaphragm reaches the bottom of
the valve, it is seated firmly, forming a leak-proof seal.
This configuration is recommended for fluids containing solid
particles and for medium-pressure service. Depending on the
diaphragm material, this type of valve can be used at
temperatures as high as 205 °C and in strong acid service. If
the diaphragm fails, the valve can become a relatively larger
source of emissions.
Diaphragm valves are very corrosion resistant and have
reportedly performed well as control valves with minimum
maintenance. The design problems associated with these valves
are the temperature and pressure limitations of the elastomer
used for the diaphragm. Use at temperatures beyond the
operating limits of the material tends to damage or destroy
the diaphragm. Also, constraints on operating pressure may
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limit the application of diaphragm valves to medium- and low-
pressure operations.
Sealed bellows valves are another alternative leakless
design for isolating the valve stem from the process fluid.
In this type of valve, metal bellows are welded to the bonnet
and disk of the valve, thereby isolating the stem from the
process. These valves can be designed to withstand high
temperatures and pressures and can provide leak-free service
at operating conditions beyond the limits of diaphragm valves.
However, they are usually dedicated to highly toxic services
and the nuclear industry. Further, they are only available in
globe and gate valve configurations, and the bellows alloy
must be carefully selected to prevent corrosion of the
crevices of the bellows under severe conditions.
The control effectiveness of both diaphragm and sealed
bellows valves is virtually 100 percent, although a failure of
the diaphragm or bellows could cause temporary emissions much
larger than those from other types of valves.
2.4.1.8 Connectors. Flanges, threaded fittings, and
other fittings used to join sections of piping and equipment
are connectors. They are used wherever pipe or other
equipment (such as vessels, pumps, valves, and heat
exchangers) require isolation or removal.
Flanges are bolted, gasket-sealed connectors. Normally,
flanges are used for pipes with diameters of 50 mm or greater
and are classified by pressure rating and face type. The
primary causes of flange leakage are poor installation and
thermal stress, which results in the deformation of the seal
between the flange faces. Flanges can become sources of
emissions when they are assembled poorly or when replacement
gaskets are improperly chosen.
Threaded fittings are made by cutting threads into the
outside end of one piece (male) and the inside end of another
piece (female). These male and female parts are then screwed
together like a nut and bolt. Threaded fittings are normally
used to connect piping and equipment having diameters of 50 mm
or less. Seals for threaded fittings are made by coating the
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male threads with a sealant (e.g., pipe dope) before joining
it to the female piece. The sealant may be a polymeric tape,
brush-on paste, or other spreadable material that acts like
glue in the joint. These sealants typically need to be
replaced each time the joint is broken. Emissions can occur
as the sealant ages and eventually cracks. Leakage can also
occur as the result of poor assembly or application of the
sealant, and thermal stress of the piping and fittings.
Emissions from connectors can be controlled by regularly
scheduled maintenance. Potential emissions can be reduced by
replacing the gasket or sealant materials. Because most
connectors cannot be isolated for maintenance or replacement
during process operation, any maintenance must occur during
shut-downs. In cases where connectors are not required for
process modification or periodic equipment removal, emissions
from connectors can be eliminated by welding the connectors
together.
2.4.1.9 Instrumentation Systems. An instrumentation
system is a group of equipment components used to condition
and convey a sample of process fluid to analyzers and
instruments for the purpose of determining process operating
conditions (e.g., composition, pressure, and flow rate).
Valves and connectors are the predominant types of equipment
used in instrumentation systems, although other equipment may
be included. For the purposes of identification, only valves
nominally 12.7 mm and smaller and connectors nominally 19 mm
and smaller are typically considered instrumentation systems.
Emissions resulting from the components in the instrumentation
system are controlled as they are for the same component in
the process system.
2.4.1.10 Product Accumulator Vessels. Product
accumulator vessels are small, primarily fixed roof storage
tanks designed to regulate material flow through a process.
They include the overhead and bottoms receiver vessels used
with fractionation columns and the product separator vessels
used in series with reactor vessels to separate reaction
products. Under normal operating conditions, the amount of
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material entering and exiting the vessel is the same and the
net flux in the vessel is zero. Emissions occur when VOC's
are vented to the atmosphere either directly or through a
blowdown drum or vacuum system. Control of emissions from
product accumulator vessels require the capture and
transportation of VOC's through a closed vent system to a
control device.
2.4.2 Closed Vent Systems and Control Devices
Emissions from equipment leaks may be controlled by
installing a closed vent system around the leaking equipment
and venting the emissions to a control device. This method of
control is only applicable to certain equipment types (e.g.,
pumps, compressors, agitators, pressure relief valves, and
product accumulator vessels). Because of the many valves,
connectors, and open-ended lines typically found in chemical
facilities, it is not practical to use this technique for
reducing emissions from all of these potential sources for an
entire process unit. However, a closed vent system can be
used to control emissions from a limited number of components,
which could be enclosed and maintained under negative pressure
and vented to a control device.
For pumps, compressors, agitators, and product
accumulator vessels, a closed vent system would require some
type of flow-inducing device to transport the emissions from
the seal area of the piece of equipment to the control device.
The seal area must be enclosed to collect the emissions, and a
vacuum eductor or a compressor could be used to remove vapors
from the seal area. However, normal operating practices for
some of this equipment may require frequent visual inspection
or mechanical adjustment in the seal area. Also, explosive
mixtures may be created by enclosing some seal areas.
Therefore, safety and operating practices limit the use of
closed vent streams in some applications.
2.4.2.1 Combustion Control Devices. Combustion control
devices that can be used to control VOC emissions collected in
closed vent systems include incinerators, industrial boilers,
process heaters, and flares. Control efficiencies of the
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various techniques depend on operating characteristics and the
properties of the VOC being controlled.
Enclosed combustion devices include incinerators,
boilers, and process heaters. All three are capable of
achieving greater than 95 percent combustion efficiency
depending on the operating parameters of the individual
combustion device and the vent stream being controlled.
Thermal incinerators, catalytic incinerators, and boilers and
process heaters are described in detail, including discussion
of their efficiencies and applicabilities, in Sections 2.1.2,
2.1.3, and 2.1.4, respectively.
Flares are commonly used in chemical plants but have
limited application for controlling the low flow rate and low
concentration streams associated with equipment leaks. Flares
designed to handle large volumes of vapors associated with
overpressure releases may also be used to handle low volumes
of equipment leak emissions. However, optimum mixing is not
achieved in these cases because the vent stream exit velocity
is low and large flares usually cannot properly inject steam
into low volume streams. Flares are described in
Section 2.1.1, which includes a discussion of flare efficiency
and applicability.
2.4.2.2 Vapor Recovery Systems. Vapor recovery systems
collect VOC without destroying them. Condensers, adsorbers,
and absorbers are all examples of vapor recovery systems and
are described in Sections 2.2.2, 2.2.4, and 2.2.5,
respectively. These systems all require careful design and
operating practices that depend on the VOC species present. A
design or operating procedure that works well for one VOC may
not work well for other compounds.
2.4.3 Work Practices
Leak detection and repair methods are used to identify
equipment components that are emitting significant amounts of
VOC and to reduce these emissions. The emission reduction
potential for LDAR as a control technique is highly variable
and depends on several factors, the most important of which
are the frequency of monitoring and the techniques used to
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identify leaks. Monthly monitoring is typically more
effective in reducing emissions than quarterly monitoring
because leaks can be identified and repaired more quickly.
Repair of leaking components is required only when the
equipment leak emissions reach a set level—the leak
definition level. A low leak definition will instigate repair
at lower levels, resulting in a lower overall emission rate.
The leak occurrence rate, leak recurrence rate, and repair
effectiveness of individual sources also affect emission
reductions.
Emissions from leaking sources can be reduced by repair,
modification, or replacement of the source. The following
sections describe various leak detection programs and repair
methods to reduce emissions from the leaking equipment.
2.4.3.1 Leak Detection Methods. Leak detection methods
include individual component surveys, area (walk-through)
surveys, and fixed point monitors. Individual component
surveys form a part of the other methods.
2.4.3.1.1 Individual component survey. Each source of
equipment leak emissions (pump, valve, compressor, etc.) can
be checked for leakage of VOC by visual, audible, olfactory,
soap bubble, or instrument techniques. Visual methods are
good for locating liquid leaks, especially large pump seal
failures. A visible leak does not necessarily indicate VOC
emissions, however, because the leaking material may be non-
VOC. High-pressure leaks may be detected by the sound of
escaping vapors, and leaks of odorous materials may be
detected by smell.
Spraying soap on equipment components can be used to
survey individual components in certain applications. If the
soap solution forms bubbles or blows away, a leak is
indicated. If bubbles do not form in the soap solution and
the soap solution is not blown away, no leak is detected.
Disadvantages of this method are that (1) it does not
distinguish leaks of hazardous VOC's from leaks of
nonhazardous VOC's; (2) it is only semiquantitative, because
it requires the observer to determine subjectively the rate of
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leakage based on behavior of the soap bubbles; and (3) it is
limited to "cool" sources, because temperatures above 100 °C
would cause the water in the soap solution to evaporate. This
method is also not suited for moving shafts on pumps or
compressors, because the motion of the shaft may interfere
with the motion of the bubbles caused by a leak.
Using a portable hydrocarbon detection instrument is the
best method for identifying leaks of VOC from equipment
components. The instrument is used to sample and analyze the
air close to the potential leak surface by traversing the
sampling probe tip over the entire area where leaks may occur.
This sampling traverse is called "monitoring." The
hydrocarbon concentration of the sampled air is displayed on
the instrument meter and is an indication of the VOC emission
rate from the component. Components that have indicated
concentrations higher than a specified "action level" are
marked for repair.
2.4.3.1.2 Area survey. An area survey (or walk-through
survey) requires the use of a portable hydrocarbon detector
and a strip chart recorder. The procedure involves carrying
the instrument within 1 m of the upwind and downwind sides of
process equipment and associated sources of equipment leak
emissions. The instrument is then used for an individual
component survey in a suspected leak area. The efficiency of
this method for locating leaks is not well established. The
time and labor requirements for the walk-through are much
lower than for an individual component monitoring survey, but
this method will not detect leaks from sources such as
overhead valves or relief valves. Leaks from adjacent units
and adverse meteorological conditions can affect the results
of the walk-through survey. Consequently, the walk-through
survey is best for locating only large leaks at small expense.
2.4.3.1.3 Fixed point monitors. This method consists of
placing several automatic hydrocarbon sampling and analysis
instruments at various locations in the process unit. The
instruments may sample the ambient air intermittently or
continuously. Elevated hydrocarbon concentrations indicate a
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leaking component. As in the walk-through method, an
individual component survey with a portable hydrocarbon
detector is required to identify the specific leaking
component in the area. Leaks from adjacent units and
meteorological conditions may affect the results. The
efficiency of this method is not well established, but fixed
point monitoring of VOC's is not as effective as a complete
individual component survey. 7 This monitoring method is
expensive because fixed-point monitors are expensive, multiple
units may be required, and the portable instrument is also
needed to locate the particular leaking component.
Calibration and maintenance costs may be high. Fixed-point
monitors are used successfully to detect emissions of
hazardous or toxic substances (such as vinyl chloride) as well
as potentially explosive conditions. Fixed-point monitors can
provide an increased detection efficiency by selecting a
particular compound as the sampling criterion.
2.4.3.2 Repair Methods. The following descriptions of
repair methods include only those features of each equipment
emission source (pumps, valves, etc.) that need to be
considered. They are not intended to be complete repair
procedures.
Many pumps have in-line or parallel spares that can be
used while the leaking pump is being repaired. Leaks from
packed seals may be reduced by tightening the packing gland.
The packing may deteriorate to the point where further
tightening has no effect or even increases emissions from the
seal. At this point, the packing can be replaced with the
pump out of service. With mechanical seals, the pump must be
dismantled to repair or replace the leaking seal. Dismantling
pumps can result in spillage of some process fluid. If the
seal leak is small, evaporative emissions of VOC from such
spillage may be greater than the continued leak from the seal.
Precautions must be taken to prevent or reduce these
emissions.
Leakage from compressors with packed seals may be reduced
by tightening the packing gland, as described for pumps.
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Repair of compressors with mechanical seals requires the
compressor be removed from service. Because compressors
typically do not have spares, immediate repair may not be
practical or possible without a process unit shutdown.
Like pumps and compressors, agitators can leak organic
chemicals at the point where the shaft penetrates the casing,
and seals are required to minimize equipment leak emissions.
Leaks from packed seals may be reduced by the repair procedure
described for pumps. Repair of other types of seals require
the agitator to be out of service, which often requires
shutdown of the process or isolation of the particular
agitator being repaired. If the leak is small, temporary
emissions resulting from a shutdown may be greater than the
emissions from the leaking seal; precautions must be taken to
prevent or reduce these emissions.
Relief valves that leak usually must be removed for
repair. In some cases of improper reseating, manual release
of the valve may improve the seat seal. To remove the relief
valve without shutting down the process, a block valve may be
required upstream of the relief valve. A spare relief valve
should be attached while the faulty valve is repaired and
tested. The repair and replacement of a relief valve does not
guarantee that the next overpressure relief will not result in
another leak.
A rupture disk can be installed upstream from a pressure
relief valve to eliminate leaks until an overpressure release
occurs. Once a release occurs, the rupture disk must be
replaced to prevent further leaks. A block valve is required
to isolate the rupture disk for replacement.
Most valves have a packing gland that can be tightened
while in service. Although this procedure should decrease the
emissions from the valve, it can actually increase the
emission rate if the packing is old and brittle or has been
over-tightened. Plug-type valves can be lubricated with
grease to reduce emissions around the plug. Some types of
valves have no means of in-service repair and must be isolated
from the process and removed for repair or replacement. Other
2-82
-------
valves, such as control valves, may be excluded from in-
service repair by operating or safety procedures. Most
control valves have a manual bypass loop that allows them to
be isolated and removed. Most block valves cannot be isolated
easily, although temporary changes in process operation may
allow isolation in some cases. The emissions resulting from
the shutdown of a process unit to isolate a leaking valve may
be greater than those occurring from the valve before the next
process change (when the valve could be more easily repaired).
Depending on site-specific factors, it may be possible to
repair process valves by injecting a sealing fluid into the
source. This type of repair may affect the operability of the
valve and necessitate replacing the valve shortly after its
repair. Injection of sealing fluid has been used successfully
to repair leaks from valves in petroleum refineries in
_ n . _ . 68
California.
In some cases, leaks from connectors can be reduced by
replacing the connector gaskets, but most connectors cannot be
isolated to permit gasket replacement. Tightening of
connector bolts also may reduce emissions from connectors.
Where connectors are not required for process modification or
periodic equipment removal, emissions from connectors can be
eliminated by welding them.
2-83
-------
2.5 REFERENCES
1. Kalcevic, V. (IT Enviroscience). Control Device
Evaluation: Flares and the Use of Emissions as Fuels.
In: Organic Chemical Manufacturing, Volume 4:
Combustion Control Devices. Report 4. Prepared for
U.S. Environmental Protection Agency. Research Triangle
Park, NC. Publication No. EPA-450/3-80-026. December
1980.
2. Klett, M.G. and J.B. Galeski (Lockheed Missiles and Space
Co., Inc.). Flare Systems Study. Prepared for U.S.
Environmental Protection Agency. Research Triangle Park,
NC. Publication No. EPA-600/2-76-079. March 1976.
3. U.S. Environmental Protection Agency, Office of Air
Quality Planning and Standards. Evaluation of the
Efficiency of Industrial Flares: Background—
Experimental Design—Facility. EPA-600/2-83-070.
Research Triangle Park, NC. August 1983.
4. U.S. Environmental Protection Agency, Office of Air
Quality Planning and Standards. Reactor Processes in
Synthetic Organic Chemical Manufacturing Industry—
Background Information for Proposed Standards.
EPA 450/3-90-016a. Research Triangle Park, NC. June
1990.
5. Letter from Matey, J.S., Chemical Manufacturers
Association, to Beck, D., EPA/CPB. November 25, 1981.
Response to questionnaire on distillation technical
issues.
6. Reed, R.J. North American Combustion Handbook.
Cleveland, OH. North American Manufacturing Company.
1978. p. 269.
7. U.S. Environmental Protection Agency, Office of Air
Quality Planning and Standards. OAQPS Control Cost
Manual. Fourth Edition. EPA-450/3-90-006. Research
Triangle Park, NC. January 1990. p. 3-43.
8. Telecon. Stone, O.K., Radian Corporation, with Dowd, E.,
ARI Technology. January 18, 1990. Incinerator sizes and
turndown.
9. Memorandum and attachments from Farmer, J.R., EPA/ESD, to
Ajax, B. et al. August 22, 1980. Thermal incinerators
and flares.
-------
10. Kosusko, M., and G. Ramsey. Destruction of Air Emissions
Using Catalytic Oxidation. U.S. Environmental Protection
Agency, Air and Energy Engineering Research Laboratory.
Research Triangle Park, NC. Publication No.
EPA/600-D-88/107. May 1988. p. 3.
11. U.S. Environmental Protection Agency, Office of Air and
Waste Management. Control Techniques for Volatile
Organic Emissions from Stationary Sources.
EPA-450/2-78-002. Research Triangle Park, NC.
May 1978. p. 32.
12. Ref. 10, p. 5.
13. Kenson, R.E. A Guide to the Control of Volatile Organic
Emissions. MetPro Corp., Systems Division.
Harleysville, PA. S104A.5. August 1981.
14. Ref. 10, p. 2.
15. Devitt, T., P. Spaite, and L. Gibbs (PEDCo Environmental,
Inc.). Population and Characteristics of
Industrial/Commercial Boilers in the U.S. Prepared for
U.S. Environmental Protection Agency. Washington, DC.
Publication No. EPA 600/7-79-178a. August 1979.
16. U.S. Environmental Protection Agency, Office of Air
Quality Planning and Standards. Fossil Fuel Fired
Industrial Boilers - Background Information, Volume 1:
Chapters 1-9. EPA-450/3-82-006a. Research Triangle
Park, NC. March 1982. p. 3-27.
17. Hunter, S.C. and S.S. Cherry (KVB). NOX Emissions from
Petroleum Industry Operations. Prepared for the American
Petroleum Institute. Washington, DC. API Publication
No. 4311. October 1979. p.83.
18. Castaldini, C., H.K. Willard, D. Wolbach, and
L. Waterland (Acurex Corporation). A Technical Overview
of the Concept of Disposing of Hazardous Wastes in
Industrial Boilers. Prepared for the U.S. Environmental
Protection Agency. Cincinnati, OH. EPA Contract No. 68-
03-2567. January 1981. p. 44.
19. Ref. 18, p. 73.
20. U.S. Environmental Protection Agency. Benzene—Organic
Chemical Manufacturing. Emission Test Report,
Ethylbenzene/Styrene, Amoco Chemicals Company (Texas
City, Texas). EMB Report No. 79-OCM-13. Research
Triangle Park, NC. August 1979.
2-85
-------
21. U.S. Environmental Protection Agency. Benzene—Organic
Chemical Manufacturing, Ethylbenzene Styrene. Emission
Test Report, El Paso Products Company (Odessa, Texas).
EMB Report No. 79-OCM-15. Research Triangle Park, NC.
April 1981.
22. U.S. Environmental Protection Agency. Benzene—Organic
Chemical Manufacturing, Ethylbenzene Styrene. Emission
Test Report, USS Chemicals (Houston, Texas). EMB Report
No. 80-OCM-19. Research Triangle Park, NC. August 1980.
23. U.S. Environmental Protection Agency, Office of Air
Quality Planning and Standards. OAQPS Control Cost
Manual. Fourth Edition. Supplement 1. Chapter 8. EPA-
450/3-90-006a. Research Triangle Park, NC. January
1992.
24. Erikson, D.G. (IT Enviroscience). Control Device
Evaluation: Condensation. In: Organic Chemical
Manufacturing, Volume 5: Adsorption, Condensation, and
Absorption Devices. Prepared for U.S. Environmental
Protection Agency. Research Triangle Park, NC.
Publication No. EPA-450/3-80-027. December 1980. Report
2. p. II-2.
25. Green, D.W., ed. Perry's Chemical Engineers' Handbook,
Sixth Edition. New York, McGraw Hill Book Company.
1984. pp. 12-29, 12-30.
26. Letter from Plant A to Farmer, J., EPA/ESD. October
1986. Confidential Section 114 response.
27. Letter from Plant B to Farmer, J., EPA/ESD. November
1986. Confidential Section 114 response.
28. Trip Report. Howie, R.H., and M. A. Vancil, Radian
Corporation, to file. 7 p. Report of May 12, 1987,
visit to Allied Fibers.
29. Memorandum from Zukor, C.J., Radian Corporation to
Lassiter, P., EPA/CPB. January 27, 1992. Development of
National Emission Impacts from Responses to the March
1990 Section 114 Wastewater Questionnaire: A Summary of
the Section 114 Database.
30. Stern, A.C. Air Pollution. Third Edition. Volume IV.
New York, Academic Press. 1977. p. 336.
31. Barnett, K.W., P.A. May, and J.A. Elliott (Radian
Corporation). Carbon Adsorption for Control of VOC
Emissions: Theory and Full Scale System Performance.
Prepared for U.S. Environmental Protection Agency.
Research Triangle Park, NC. EPA Contract No. 68-02-4378.
June 6, 1988. p. 2-2.
-------
32. Basdekis, H.S., and C.S. Parmele (IT Enviroscience).
Control Device Evaluation: Carbon Adsorption. In:
Organic Chemical Manufacturing, Volume 5: Adsorption,
Condensation, and Absorption Devices. Report 1. U.S.
Environmental Protection Agency. Research Triangle Park,
NC. Publication No. EPA-450/3-80-027. December 1980. p.
II-l.
33. Ref. 32, p. II-7
34. Ref. 31, p. 3-16.
35. U.S. Environmental Protection Agency, Air and Energy
Engineering Research Laboratory. Handbook—Control
Technologies for Hazardous Air Pollutants. Publication
No. EPA-625/6-86-014. Research Triangle Park, NC.
September 1986. p. 26-27.
36. Research and Education Association. Modern Pollution
Control Technology, Volume I. New York. 1978. pp. 22-
23.
37.
38.
39.
40.
41.
42.
43.
44.
Ref.
Ref.
Ref.
Ref.
Ref.
Ref.
Ref.
31,
35,
30,
11,
24,
25,
11,
Standif er
pp.
P-
P-
P-
P-
P-
P-
, **
3-34 and 3-38.
65.
24.
72.
II-l.
14-1.
76.
:.L. (IT Enviros
Evaluation: Gas Absorption. In: Organic Chemical
Manufacturing, Volume 5: Adsorption, Condensation, and
Absorption Devices. Report 3. U.S. Environmental
Protection Agency. Research Triangle Park, NC.
Publication No. EPA-450/3-80-027. December 1980.
p. 1-1.
45. Ref. 44, p. 1-7.
46. Memorandum from Probert, J.A., Radian Corporation, to HON
project file. August 7, 1991. Achievable emission
reduction for internal floating roofs.
47. Memorandum from Zukor, C.J., Radian Corporation, to HON
project file. July 29, 1991. Emission rates from
example internal floating roof tanks with various
combinations of rim seals, deck fittings, and deck seams.
2-87
-------
48. U.S. Environmental Protection Agency, Office of Air
Quality Planning and Standards. VOC Emissions from
Volatile Organic Liquid Storage Tanks—Background
Information for Proposed Standards. EPA-450/3-81-003a.
Research Triangle Park, NC. p. 4-15. July 1984.
49. Ref. 48, Appendix C.
50. Letter and attachments from Blumquist, F.L., PETREX,
Inc., to McDonald, R.J., EPA/CPB. April 26, 1991.
Submittal of test data on gasketing deck seams.
51. U.S. Environmental Protection Agency. Fugitive Emission
Sources of Organic Compounds—Additional Information on
Emissions, Emission Reductions, and Costs. Section 5.
EPA-450/3-82-010. Research Triangle Park, NC.
April 1982.
52. Erikson, D. G. and V. Kalcevic. (IT Enviroscience, Inc.)
Fugitive Emissions. In: Organic Chemical Manufacturing,
Volume 3: Storage, Fugitive, and Secondary Sources.
Report 2. U.S. Environmental Protection Agency.
Research Triangle Park, NC. EPA-450/3-80-025. December
1980. p. II-2.
53. Ramsden, J. H. How to Choose and Install Mechanical
Seals. Chem. Eng. 85(22):97-102. October 9, 1978.
54. Ref. 53, p. 99.
55. Ref. 52, p. IV-I.
56. Nelson, W. E. Compressor Seal Fundamentals. Hydrocarbon
Pro. 56(12):91-95. December 1977.
57. Ref. 52, p. II-6.
58. Ramsey, W. D. and G. C. Zoller. How the Design of
Shafts, Seals, and Impellers Affects Agitator
Performance. Chem. Eng. 83 (18): 101-108. August 30,
1976.
59. Ref. 58, p. 105.
60. American Petroleum Institute. Guide for Pressure Relief
and Depressuring Systems. API RP 521. Washington, DC.
September 1969.
61. Ref. 52, p. 11-13.
62. Lyons, J. L. and C. L. Askland. Lyons' Encyclopedia of
Valves. New York, Van Nostrand Reinhold Company. 1975.
290 p.
2-88
-------
63. Templeton, H. C. Valve Installation, Operation and
Maintenance. Chem. Eng. 78(23):141-149. October 11,
1971.
64. Ref. 63, p. 147-148.
65. Pikulik, A. Manually Operated Valves. Chem. Eng.
85(7):121. April 3, 1978.
66. McFarland, I. Preventing Flange Fires. Chem. Eng Progr.
65(8):59-61. August 1969.
67. Hustvedt, K. C. and R. C. Weber. (U. S. Environmental
Protection Agency.) Detection of Volatile Organic
Compound Emissions from Equipment Leaks. Presented at
the 71st Annual Meeting of the Air Pollution Control
Association. Houston, TX. June 25-30, 1978.
68. Teller, J. H. Advantages Found in On-Line Leak Sealing.
Oil Gas J. 77(29): 54-59. July 16, 1979.
2-89
-------
3.0 COST ANALYSIS
This chapter presents methodologies for costing selected
technologies described in Chapter 2 for control of VOC
emissions from the five source types in the organic chemical
manufacturing industry. The presentation of these methods
includes discussion of the design aspects of these
technologies that influence costs.
3.1 COST METHODOLOGY FOR COMBUSTION SYSTEMS
This section presents the costing methodologies for two
control technologies that reduce VOC emissions by destruction
of the organic compounds through combustion. Equipment
associated with these systems, design assumptions, and costing
equations are provided for flares and incinerators. The
incinerator cost includes a packed tower scrubber system that
removes acidic vapors from the incinerator flue gas for
streams containing halogenated VOC's.
3.1.1 Cost Methodology for Flare Systems
Flare design and costs are based on Chapter 7 of the
OAQPS Control Cost Manual (OCCM).1
3.1.1.1 Flare Design Considerations Affecting Cost. The
flare costed in this section is an elevated, steam-assisted,
smokeless flare. Elements of the flare system include the
knockout drum, liquid seal, stack, gas seal, burner tip, pilot
burners, and steam jets. To size the flare system,
correlations were developed relating process vent stream flow
rate and heat content to the flare height and tip diameter.
The general design specifications used in developing these
correlations are presented in Table 3-1.
Flare height and tip diameter are the primary design
parameters used to determine the capital cost of a flare. The
tip diameter selected is a function of the combined flow rates
3-1
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of the vent stream and supplemental fuel, and the assumed tip
velocity. Supplemental fuel requirements and tip velocity
values are listed in Table 3-1. The maximum allowed velocity
ranges from 18.3 to 121.9 m/sec (60 to 400 ft/sec), as
required by rules defined in the Federal Register. The flare
tip diameter is sized to allow a flare tip velocity of
80 percent of the maximum design velocity. Determination of
flare height is based on worker safety requirements and is
selected so the maximum heat intensity at ground level
including solar radiation is 1,576 W/m2 (500 Btu/hr ft2).
Vendor contacts indicate the smallest elevated flare
commercially available is 2.54 cm (I in.) in diameter and
9.14 m (30 ft) high.
After the flare tip diameter, D, and flare height, H, are
determined, the natural gas required for pilots and purge and
the mass flow rate of steam required are calculated. Pilot
gas consumption is a function of the number of pilots, which
is selected based on the tip diameter (Table 3-1). Pilot gas
consumption is calculated based on an energy-efficient model
of 2 scm/hr (70 scf/hr) per pilot burner. The purge gas
requirement is also a function of the tip diameter and the
minimum design purge gas velocity of 0.012 m/sec (0.04 ft/sec)
at the tip (Table 3-1). Steam is required to maintain a
steam-to-flare-gas ratio of 0.4 kg steam/kg flare gas.
3.1.1.2 Development of Flare Capital Costs. The capital
cost of a flare is based on vendor-supplied information as
correlated in the OCCM cost equations, regressed from the
combined data set over a range of tip diameters and flare
heights. Flare equipment costs, CF, are calculated based on
stack height, H (ft), and tip diameter, D (in.), according to
type of support for the flare:
• Self supported:
CF (July 1989 dollars) = [78.0 + 9.14(D) +
0.749(H)]2
• Guy supported:
CF (July 1989 dollars) = [103 + 8.68(D) + 0.470(H)]2
3-3
-------
• Derrick supported:
CF (July 1989 dollars) = [76.4 + 2.72(D) + 1.64(H)]2
The total cost for flare equipment includes the costs of the
flare tower (stack) and support, burner tip, pilots, piping
for utilities, piping from base, meters and controls for
utilities, liquid seal, gas seal, and galvanized caged ladders
and platforms as required. The material of construction is
carbon steel, except for the upper 4 ft and burner tip, which
is 310 stainless steel.
Vent stream piping costs are a function of flare tip
diameter and length of piping. The pipe diameter is assumed
to be equal to the flare tip diameter. Pipe costs are
calculated using the following equation:
Cp (July 1989 dollars) = 1.27 * D1-21 * 400 for D < 24
Cp (July 1989 dollars) = 1.39 * D1-07 * 400 for D > 24
where:
Cp = vent stream piping costs; and
D = diameter (in.)
These costs include 400 ft of Schedule 40, carbon steel pipe.
Knockout drum costs are a function of drum diameter (in.)
and drum wall thickness (in.):
CK = 14.2 {(d)*(t)*[h + 0.812*(d)]}0.737
where:
CK = knockout drum costs;
d = drum diameter; and
t = drum wall thickness
The cost of collection fan, Cpan, is a function of total
stream flow rate, Qtot (scfm):
Cfan = 96.96 * Qtot °'5472
Total flare system equipment cost is the sum of flare,
piping, knockout drum, and fan costs:
EC = CF + CK + Cp + cFan
where:
EC = equipment costs;
Cp = flare cost;
CK = knockout drum cost; and
Cp = piping costs.
3-4
-------
Purchased equipment cost equals equipment cost plus factors
for auxiliary equipment (i.e., instrumentation at 0.10, sales
tax at 0.03, and freight at 0.05). Installation cost is
estimated as a percentage of purchased equipment cost. The
total capital investment is obtained by multiplying the
purchased equipment cost by an installation factor of 1.92.
Further details about flare costs are presented in Chapter 7
of the OCCM.1
3.1.1.3 Development of Flare Total Annual Cost. The
annual costs for a flare system include direct operating and
maintenance costs, and annualized capital charges. The
assumptions used to determine annual costs are presented in
Table 3-2, and are given in July 1989 dollars. Direct
operating and maintenance costs include operating and
maintenance labor, maintenance materials, and utilities.
3.1.1.3.1 Labor costs. The operating labor requirements
are 630 hr/yr for typical flare systems. Maintenance labor
requirements are .5 hr per 8-hr shift. Supervisory labor is
estimated to be 15 percent of the operating labor cost. The
hourly rate for maintenance labor is assumed to be 10 percent
higher than that for operating labor.
3.1.1.3.2 Capital charges. Capital recovery costs are
estimated using the following equation:
CRC = CRF * TCI
where:
CRC = capital recovery charges;
CRF = capital recovery factor; and
TCI = total capital investment.
The capital recovery factor (0.1314) is based on a 10 percent
interest rate and a 15-year life for the equipment. Taxes,
insurance, and administrative costs are assumed to be
4 percent of the total capital investment. Also, overhead
costs are assumed to be 60 percent of all labor and
maintenance costs.
3.1.1.3.3 Utility costs. The utilities considered in
the annual cost estimates include steam, natural gas, and
electricity. Cost of steam to eliminate smoking is equal to
3-5
-------
TABLE 3-2.
BASES AND FACTORS FOR ANNUAL COSTS
FOR FLARES
Basis for Direct Annual Costs
Operating labor (man-hours/8-hr shift)
Maintenance labor (man-hours/8-hr shift)
Labor rates (July 1989 $/hr)a
Operating labor
Maintenance labor
Supervisory labor cost
(percent of operating labor cost)
Maintenance materials cost
(percent of maintenance labor cost)
Utilities (July 1989 $)a
Electricity ($/l,000 KW-hr)
Natural Gas ($/1,000 ft3)
Steam ($/Mg)
Basis for Indirect Annual Cost
Equipment life (years)
Interest rate (percent)
Capital recovery factor
Taxes, insurance, administration (percent of
total capital investment)
Overhead (percent of total
labor and maintenance costs)
Factor
0.575
0.5
13.20
14.50
15
100
50.9
3.03
7.68
15
10
0.1314
4
60
aBased on Reference 3.
3-6
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the annual steam consumption multiplied by the unit cost.
Electrical costs for operating a fan are proportional to total
stream flow rate. Flares use natural gas in three ways: for
pilot burners, as an auxiliary fuel for low Btu vent streams,
and as a purge gas. The procedures for estimating
electricity, steam, and supplemental fuel requirements are
described in Section 3.1.1.1 and in Chapter 7 of the OCCM.1
3.1.1.3.4 Maintenance costs. Maintenance material costs
are assumed to equal maintenance labor costs, which are
discussed in Section 3.1.1.3.1.
3.1.2 Cost Methodology for Incinerator Systems
Incinerator costs were developed using Chapter 3 of the
OCCM. Scrubber costs were developed using the procedure
outlined in EPA's Handbook on Control Technologies for
Hazardous Air Pollutants (HAP Manual),4 with equipment costs
updated from recent technical journal information. '
3.1.2.1 Thermal Incinerator Design Considerations
Affecting Cost. The thermal incinerator system consists of
the following equipment: combustion chamber, instrumentation,
recuperative heat exchanger, blower, collection fan and
ductwork, quench/scrubber system (if applicable), and stack.
The OCCM contains further discussion of incinerator control
system design. Control system elements and design assumptions
specific to SOCMI vent streams are discussed below. General
incinerator design specifications are presented in
Table 3-3.
3.1.2.1.1 Combustion air requirements. The amount of
oxygen in the vent stream or bound in the VOC establishes the
supplemental combustion air requirement, which has an impact
on both the capital and operating costs of the thermal
incinerator. Supplemental combustion air may be required to
support combustion. This cost analysis assumes that the vent
stream is a mixture of VOC, oxygen, and an inert gas such as
nitrogen. After combustion, the excess oxygen content in the
incinerator flue gas is designed to be at least 3 mole
percent, which is based on commonly accepted operating
practice.
3-7
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TABLE 3-3. INCINERATOR GENERAL DESIGN SPECIFICATIONS
Item
Specification
Emission control efficiency
Minimum incinerator capacity3
Maximum incinerator capacity
Incinerator temperature
- Nonhalogenated vent streams
- Halogenated vent streams*3
Chamber residence times
- Nonhalogenated vent streams
- Halogenated vent streams*3
Supplemental fuel requirement
Scrubber system
98 percent destruction
500 scfm
50,000 scfm
870 °C (1,600 °F)
1,100 °C (2,000 °F)
0.75 sec
1.00 sec
Natural gas required to
maintain incinerator
temperature with 3 mole
percent excess oxygen in
flue gas
Used when halogenated VOC is
present to remove corrosive
combustion by-products
Type
Packing type
Scrubbing liquid
Scrubber gas temperature
Packed tower
2-in. rings, ceramic
Water
25 °C (77 °F)
aFive hundred scfm is the minimum incinerator s^.ze used to
determine capital cost. A 10:1 turndown ratio is assumed
and a minimum flow rate of 50 scfm is used to determine
operating costs.
bUsed when halogenated VOC are present due to the difficulty
of achieving complete combustion of halogenated VOC at lower
temperatures.
3-8
-------
To calculate the amount of combustion air required to
ensure a flue gas ©2 concentration of 3 mole percent, a
complete stoichiometric equation must be balanced for each
compound present in the vent stream. In many cases, the
complete chemical composition of the vent stream is not known.
Thus, to cost incinerator systems for typical vent streams
encountered, a design molecule approach was used. The design
molecule was based on a survey of typical values for carbon,
hydrogen, oxygen, sulfur, and chloride ratios for a group of
o p
219 organic compounds. ' For the analysis, the average VOC
molecular composition of 68.3 percent carbon, 11.4 percent
hydrogen, and 20.3 percent oxygen was used to calculate
combustion air requirements. These weight ratios correspond
to a molecular formula of 02.85^5.700.63- If zero percent 02
in the waste stream is assumed, a dilution ratio (mole of air
per mole of VOC) of approximately 18:1 is required to achieve
3 percent ©2 in the incinerator flue gas. If sufficient
oxygen is present to ensure 3 percent ©2 in the flue gas, no
combustion air is added.
3.1.2.1.2 Dilution air requirements. After the required
combustion air is calculated and added to the total vent
stream flow, the overall heat content (Btu/scf) of the stream
is recalculated. Adding combustion air will effectively
dilute the stream and lower the heat content of the combined
stream fed to the incinerator. However, if the heat content
of the vent stream is still greater than 98 Btu/scf for
nonhalogenated streams or greater than 95 Btu/scf for
halogenated streams, additional dilution air must be added to
reduce the heat content to these acceptable levels. The
imposition of a maximum heat content level prevents the
temperature in the incinerator from exceeding the design
specifications.
The minimum flow rate to the incinerator is 50 scfm. It
is assumed that vent streams smaller than 50 scfm will be
mixed with air to achieve this minimum flow rate. The maximum
incinerator flow rate is 50,000 scfm. Flow rates greater than
3-9
-------
50,000 scfm are assumed to be handled by multiple
incinerators.
3.1.2.1.3 Recuperative heat recovery. Halogenated vent
streams are not considered candidates for heat recovery
systems and are costed assuming zero percent heat recovery.
This conservative design assumption is imposed because of the
potential for corrosion in the heat exchanger and incinerator.
If the temperature of the flue gas leaving the heat exchanger,
Tf0, were to drop below the acid dew point temperature, acid
gases would condense. Significant corrosion can lead to
shortened equipment life, higher maintenance costs, and
potentially unsafe working conditions.
Nonhalogenated vent streams are considered candidates for
recuperative heat recovery. The extent of heat recovery
depends on the .ieat content of the vent stream after dilution.
Four different heat recovery scenarios are evaluated for
nonhalogenated streams. The cost algorithm includes systems
with 0, 35, 50, and 70 percent heat recovery. The extent of
heat exchange to be used is decided by an economic
optimization procedure with the following restrictions. No
heat recovery is allowed for vent streams with a heat value
greater than 25 percent of the LEL because of the possibility
of explosion or damaging temperature excursions within the
heat exchanger. This limit typically corresponds to a heat
content of 13 Btu/scf. Therefore, if the heat content of the
total vent stream is still greater than 13 Btu/scf, even after
addition of required combustion and dilution air, then the
vent stream cannot be preheated. For streams with a heat
content less than 13 Btu/scf, the entire stream is preheated
in the recuperative heat exchanger, thus allowing for maximum
energy recovery. However, for streams with a heat content
greater than 13 Btu/scf, only the combustion/dilution air
stream can be preheated. In this case, the cost optimization
procedure evaluates the option of preheating only the air
stream and combines the VOC stream with the preheated air
stream in the incinerator.
3-10
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All allowable heat recovery percentages are evaluated,
and the calculated total capital and annual costs are based on
the most cost effective configuration. The trade-off between
the capital cost of the equipment and the operating cost
(fuel) of the system determines the optimum level of energy
recovery.
3.1.2.1.4 Incinerator design temperature. The
destruction ,of VOC is a function of incinerator temperature
and residence time in the combustion chamber. The design VOC
destruction efficiency is 98 percent, which can be met by
well-designed and well-operated thermal incinerator systems.
Previous EPA studies show that 98 percent destruction
efficiency can be met in a thermal incinerator operated at a
temperature, Tf^, of 1600 °F and a residence time of
0.75 sec.1 Thermal oxidation of halogenated VOC requires
higher temperature oxidation to convert the combustion product
to a form that can be more readily removed by flue gas
scrubbing. For instance, chloride-containing vent streams are
burned at high temperature to convert the chlorine to HC1
instead of to Cl2, because HC1 is more easily scrubbed.
Available data indicate that a temperature of 2,000 °F and
residence time of 1 sec are necessary to achieve 98 percent
VOC destruction efficiency for halogen-containing vent
streams .
3.1.2.1.5 Acid gas scrubber design consideration. Acid
gas scrubbers are packed towers designed to promote absorption
of acidic combustion by-products. Scrubbers consist of the
following major equipment: quench chamber, absorption column
and platform, packing, piping, ductwork and blowers, and the
release stack. The HAP Manual contains a full discussion of
scrubber design procedures. System elements and design
assumptions specific to this analysis are discussed below.
General scrubber design specifications are presented in
Table 3-4. Column diameter and height are the primary design
parameters in determining the capital cost of the scrubber.
These design parameters establish the column shell geometry
and amount of packing required.
3-11
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TABLE 3-4. SCRUBBER GENERAL DESIGN SPECIFICATIONS
Item
Specification
Emission control efficiency3
Column type3
Quench chamber temperature3
Scrubbing liquid3
Packing type3
Packing constant*3/c:
a
b
c
d
e
s
Y
g
r
Absorption factor3
Schmidt number for acid streams (gas)d:
Hydrogen chloride
Hydrogen bromide
Hydrogen fluoride
Schmidt number for acid streams
liquid):
Hydrogen chloride*3
Hydrogen bromide^
Hydrogen fluorided
98 percent recovery
Packed tower
25 °C
Water
2-in. ceramic rings
28
3.82
0.41
0.45
0.74
0.22
0.0125
11.13
0.00295
19
0.809
1.175
0.78
381
517.1
300.6
3From Reference 4.
bFrom Reference 11.
GFrom Reference 12.
dFrom Reference 13.
12
3-12
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The design procedure assumes no heat effects are
associated with the absorption process and that both the gas
and liquid streams are dilute solutions. The system flow
rates (gas and liquid) are assumed to be constant across the
scrubber. Vendor data for suggested liquid-to-vapor flow rate
ratios throughout scrubber columns were used to develop an
average of 17 gpm of liquid per 1000 scfm of vapor. The
equilibrium curve can be approximated as a straight line with
a worst case slope of 0.1. This value reflects the highly
soluble nature of hydrogen chloride, hydrogen fluoride, and
hydrogen bromide, which results in a large driving force for
absorption.
Column diameter determination. The column diameter is
estimated based on the characteristics of the vent stream, the
absorption liquid, the packing material, and an assumed column
flooding condition. The column operating range is assumed to
be 60 percent of the flooding rate. Typically, design manuals
obtain the diameter of the column from graphical information
presented as a correlation curve for flooding rate in randomly
packed absorption towers. The abscissa for the column
diameter curves is defined by the following expression:
abscissa = (L/G)(pG/pL)°-5
where:
L = liquid mass flow rate (Ib/hr);
G = gas stream mass flow rate (Ib/hr);
PQ = gas stream density (Ib/ft^); and
PL = liquid stream density (lb/ft3).
After substituting the appropriate parameters, the abscissa
value can be determined. Graphically, estimating the column
diameter would require locating the abscissa point and
proceeding up to the flooding line and selecting the
corresponding ordinate. The ordinate is defined as:
ordinate = (Garea) 2 (a/e3) (JILO- 2) / (gcpLpG)
where:
Garea = 9as stream flux rate at flooding
conditions (lb/ft2 sec);
a,e = packing constants;
3-13
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9c = gravitational constant (ft/sec2). and
HL = viscosity of solvent (centipoise).
The flux rate is estimated by simple substitution of the
system constants, and the column diameter is calculated based
on the estimated flux rate.
To solve for Garea algebraically, the flooding line is
approximated mathematically. The abscissa is calculated as
described above. The ordinate value can be approximated by
the following equation:
ordinate = -0.9809 * (ABS)A[-0.0065*LN(ABS)] + ABSA -
0.0219
where:
ABS = abscissa.
The flux rate is then estimated by the expression:
Garea = 0.6 * [ordinate * pG * pL * gc / (a/63) * jiL0.2]0.5
This approximation method provides a reasonable estimate of
the column flux rate without performing graphical
interpretations.
Once the flux rate has been estimated, the column cross-
sectional area and corresponding diameter are determined. The
column area and diameter are calculated by:
Acolumn = MW * g/(3600 * Garea)
Dcolumn = 1-13 (AColumn)0-5
where:
Acolumn = column area (ft2);
g = gas molar flow (Ib-moles/hr);
MW = Molecular weight of entire gas
stream; and
Dcolumn = column diameter (ft).
Column height determination. The column must be tall
enough to ensure the required removal efficiency. The height
of a packed column is calculated by determining the required
number of theoretical transfer units and multiplying by the
height of a transfer unit. A transfer unit is a measure of
the difficulty of the mass transfer operation and is a
function of the solubility and concentration of the solute in
-------
the gas and liquid streams. The number of transfer units are
expressed as NOG (number of gas transfer units) or NOL (number
of liquid transfer units) , depending on whether the gas film
or liquid film resistance controls the absorption rate.
In emission control applications, gas film resistance
typically is controlling; therefore, NOG will be used in the
following calculations. The expression for packing height is:
Htpacking = NOG * HOG
where :
Htpacking = packing height (ft);
NOG = number of gas transfer units; and
HOG = height of an overall gas transfer
unit (ft) .
Although determining the number of gas transfer units is not
usually complicated when dilute solutions are involved, NOG
can be calculated using the following equation:
NOG = ln^ (VOCe/VOCo) x Fl-fl/AF)"! + fl/AFU
[1-(1/AF)]
where :
VOCe = VOC inlet concentration (ppmv) ;
VOCo = VOC exit concentration from column (ppmv) ;
and
AF = VOC absorption factor.
To calculate VOCo, the inlet VOC concentration must be
calculated and then reduced by the scrubber control
efficiency. The absorption factor used in this analysis is
19.16, based upon typical absorption conditions for strong
acids.
The calculation of the height of an overall transfer unit
is performed based upon the following equation:
HOG = HG + (1/AF) * HL
where :
HOG = height of an overall transfer unit (ft) ;
= height of a gas transfer unit (ft) ; and
= height of a liquid transfer unit (ft) .
Generalized correlations are available to calculate H and
3-15
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these are based on the type of packing and the gas and solvent
flow rates. The correlations for HG and HL are as follows:
HG = [b*(3,600 Garea)C /(L")d]*(ScG)0.5
HL = Y*(WliL)S*(ScL)0.5
where:
b,c,d,Y, and s = packing constants;
L" = liquid flow rate (lb/hr-ft2);
tiL = liquid viscosity (lb/hr-ft) ;
ScG = 9as stream Schmidt number; and
SCL = liquid stream Schmidt number.
In this calculation, the effect of temperature on the Schmidt
numbers is assumed to be negligible. The value for the
variable L" in this equation is defined as follows:
L" = L/Acolumn
Once the height of the column is determined, the total system
height is calculated based upon the following expression:
Httotal = Htpacking + 2 + 0.25 * Dcolumn
Estimation of the total system height provides the remaining
design factor required for costing purposes.
3.1.2.2 Development of Thermal Incinerator and Scrubber
Capital Costs. The costing analysis follows the methodology
outlined in the OCCM.1 Equipment cost correlations are based
on data provided by various vendors; each correlation is valid
for incinerators in the 500 to 50,000 scfm range. Thus, the
smallest incinerator size used for determining equipment costs
is 500 scfm; for flow rates above 50,000 scfm, additional
incinerators are costed.
Equipment costs for thermal incinerators are given as a
function of total volumetric throughput. Four equations were
used in the costing analysis, each pertaining to a different
level of heat recovery. After converting to July 1989
dollars, the equations are:
EC = 10772 Qtot0'2355 HR = 0 percent
EC = 13760 Qtot0.2609 HR - 35 percent
EC = 17848 Qtot0*2502 HR = 50 percent
EC = 22333 Qtot0'2500 HR = 70 percent
3-16
-------
where:
EC = equipment costs (July 1989 dollars)
Qtot - total volumetric throughput (scfm); and
HR = heat recovery.
The cost of ductwork (not included in equipment costs) is
based on a 24-in. diameter duct made of 1/8-in. thick carbon
steel with two elbows per 100 ft. The following equation is
that used by Vatavuk:6
Pipe Cost = [(210 * 24°-839) + (2 * 4.52 * 241-43)] * 3
The length of duct is assumed to be 300 ft. Collection fan
costs are calculated using methods developed by Richardson
Engineering Services and are discussed in Section 3.2.12. The
duct and fan costs are added to the equipment costs.
Purchased equipment cost is calculated by accounting for
instrumentation, taxes, and freight, which are assumed to be
18 percent of total equipment cost (including auxiliary
equipment).
Installation costs are estimated as a percentage of
purchased equipment costs. Table 3-5 lists the values of
direct and indirect installation factors for thermal
incinerators. Installation costs can be as high as 61 percent
of the purchased equipment cost. However, the OCCM suggests
that installation costs would be approximately 20 to
25 percent of the purchased equipment cost for incinerator
units handling flow rates less than 20,000 scfm.
The costing methodology for the scrubber follows the
procedure outlined in the HAP manual. Equipment costs for
scrubbers are given as a function of column weight, WTcoi, in
pounds. The following were used in the costing analysis:
EC =
1.900604 *
wtcol
1000
0.9389
* 1000 *
CE index
298.2
where:
wtcol
wtcol
(48 * Dcol * HTcol) + 39 * (Dcol)2
weight of the column (Ib);
3-17
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TABLE 3-5. CAPITAL COST FACTORS FOR THERMAL INCINERATORS3
Cost Item Factor
Direct Costs
Purchased equipment costs
Incinerator + auxiliary equipment13 As estimated, Ac
Instrumentationd 0.10 A
Sales taxes 0.03 A
Freight 0.05 A
Purchased equipment cost B = 1.18 A
Direct installation costs
Foundations and supports 0.08 B
Handling and erection 0.14 B
Electrical 0.04 B
Piping 0.02 B
Insulation for ductwork 0.01 B
Painting 0. 01 B
Direct installation cost 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 cost (1C) 0.31 B
Total capital investment = DC + 1C 1.61 B + SP + Bldg.
aFrom Reference 1.
^Ductwork and any other equipment normally not included with
unit furnished by incinerator vendor.
GThe factor "A" is the total equipment cost for the
incinerator and auxiliary equipment.
^Instrumentation controls often furnished with the
incinerator, and thus often included in the equipment costs.
3-18
-------
Dcoi = diameter of the column (ft); and
HTcol = height of the column (ft).
3.1.2.3 Development of Thermal Incinerator Total Annual
Cost. Annual costs for the thermal incinerator system include
direct operating and maintenance costs, as well as annualized
capital charges. An assumed incinerator turn-down ratio of
10:1 was used in this cost analysis. Consequently, the
minimum flow rate for determining operating costs is assumed
to be 50 scfm. Additional dilution air is added where
necessary to raise the total flow rate of the fuel, vent
stream, and air to 50 scfm. The bases for determining thermal
incinerator annual costs are presented in Table 3-6. Each
cost parameter is reviewed below.
3.1.2.3.1 Labor costs. The operating labor requirements
vary depending on the components of the overall system.
Continuously operating incinerator systems without a scrubber
require the least amount of operating labor (548 hr/yr or
0.5 hr per 8-hr shift). Systems employing a scrubber require
an additional 548 hr/yr operating labor. Maintenance labor
requirements are assumed to be identical to operating labor
requirements, that is, 548 hr/yr for the incinerator and
548 hr/yr for the scrubber. Supervisory cost is estimated to
be 15 percent of the operating labor cost. The maintenance
labor hourly rate is assumed to be 10 percent higher than the
operating labor hourly rate.
3.1.2.3.2 Capital recovery charges. The capital
recovery factor (0.163) is based on a 10 percent interest rate
and a 10-year life for the equipment. Taxes, insurance, and
administrative costs are assumed to be 4 percent of the total
capital investment. Overhead is estimated to be 60 percent of
the total labor and maintenance costs.
3.1.2.3.3 Utility costs. The utilities considered in
the annual cost estimates include natural gas and electricity.
The procedures for estimating electricity and supplemental
fuel requirements are described in Chapter 3 of the OCCM.
3-19
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TABLE 3-6. BASES AND FACTORS FOR ANNUAL COSTS
FOR THERMAL INCINERATORS
Basis for Direct Annual Costs Factor
Operating labor (man-hours/8-hr shift) 0.5
Maintenance labor (man-hours/8-hr shift) 0.5
Labor rates (July 1989 $/hr)a
Operating labor 13.20
Maintenance labor 14.50
Supervisory labor cost 15
(percent of operating labor cost)
Maintenance materials cost 100
(percent of maintenance labor cost)
Utilities (July 1989 $)a
Electricity ($/l,000 KW-hr) 50.9
Natural Gas ($/l,000 ft3) 3.03
Basis for Indirect Annual Cost
Equipment life (years) 10
Interest rate (percent) 10
Capital recovery factor 0.163
Taxes, insurance, administration (percent of 4
total capital investment)
Overhead (percent of total 60
labor and maintenance costs)
aBased on Reference 3.
3-20
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3.1.2.3.4 Maintenance costs. Maintenance material costs
are assumed to be equal to maintenance labor costs, which are
discussed in Section 3.1.2.3.1.
3.1.2.4 Scrubber Annual Cost. Annual costs for the acid
gas scrubber include such direct operating costs as scrubber
utility costs, operating labor, maintenance labor and
materials, supervisory labor, and capital recovery charges.
All labor costs are calculated using the same methodology as
that used for thermal incinerators (Table 3-6). Utility costs
include electricity and water costs.
Capital recovery charges are estimated based on a 10-year
equipment life and an interest rate for depreciation of
10 percent. Taxes, insurance, and administrative costs are
assumed to be 4 percent of the total capital investment.
Overhead is estimated as 60 percent of the total labor and
maintenance costs.
3.2 COST METHODOLOGY FOR COLLECTION SYSTEMS AND RECOVERY
DEVICES
3.2.1 Cost Methodology for Vapor Collection Systems for
Loading Racks
This section discusses the capital and annual costs for
retrofitting a transfer loading rack with a vapor collection
system, retrofitting tank trucks and tank cars to be vapor
tight and compatible with loading rack vapor collection
systems, and incorporating a nitrogen blanketing system for
the rack and vehicles. The costs were obtained from technical
work performed in support of the New Source Performance
Standards (NSPS) for bulk gasoline terminals and the NESHAP
for benzene emissions from benzene transfer operations.
3.2.1.1 Design Considerations Affecting Costs. Several
factors affect the design and the costs of retrofitting the
loading racks and vehicles. These parameters include the
following:
• Transfer loading throughput;
• Chemical vapor pressure;
• Transport vehicle type;
• Transport vehicle fleet size;
3-21
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• Transport vehicle vapor tightness; and
• Nitrogen blanketing.
The loading throughput affects the size of the loading
rack and the costs to retrofit the rack. The loading
throughput also affects the number of transport vehicles in a
fleet, which, in turn, affects the costs to make the vehicles
vapor tight and the costs to retrofit the fleet with the
proper vapor collection equipment.
The vapor pressure of a compound affects the level of
control necessary to reduce product loss. Transferring a
compound with a low vapor pressure might not require a vapor
collection system because only a small quantity of the
material would actually be vaporized and emitted to the
atmosphere.
A nitrogen blanketing system may be used to reduce the
potential for VOC vapors to accumulate in concentrations
approaching the explosive range. Nitrogen blanketing may also
be used during transfer of some organic compounds to eliminate
the moisture in the ambient air. This is usually done to
ensure that extremely hydrophilic compounds remain moisture-
free during transfer.
3.2.1.2 Development of Capital Costs. The total capital
investment includes the costs for vapor collection equipment,
retrofitting the loading rack, retrofitting the vehicle fleet
for vapor collection and vapor tightness, nitrogen blanketing,
and installation.
Vapor collection equipment includes piping, fans, and
motors. The amount of piping needed is a function of the
distance between the transfer rack and the control device. It
is assumed that the cost of this piping will be included in
the cost of the control device.
A fan is used to transport the VOC vapors from the
loading rack to the control device. The capital cost of the
fan and motor may be calculated from the following equation:
Fan Costs
(Including Motor) = 100.79 * (Flow Rate, cfm)0-5472
(July 1989 $)
3-22
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Based on information provided in the BID for the bulk
gasoline terminals NSPS, the installed cost in July 1989
dollars for converting transport vehicles from top splash
loading to bottom loading is $8,080 per tank truck and
$15,700 per tank car. The installed cost for converting a
loading rack from top loading to bottom loading is
$252,400 per rack. These values include the cost of equipment
for connecting the vehicles and racks to a vapor recovery
17 18
system. ' Approximately 97 percent of the SOCMI has
already converted its vehicles and, where necessary, loading
racks for submerged fill or bottom loading. Such facilities
will not incur the retrofit costs described above.
3.2.1.3 Development of Annual Costs. Direct annual cost
includes the electricity to run the vapor collection system
fan and the toll for the N2 blanketing system. The
electricity cost for the motor is a function of the vent
stream flow rate, the hours of operation, and the pressure
drop through the piping from the loading rack to the control
device. Because transfer operations are discrete events, not
continuous, the fan will not operate continuously. However,
it will operate during loading and for a period after each
loading to clear the lines of VOC vapors from residual liquid
in the lines. As a conservative estimate, it is assumed that
the fan will operate for a period twice that of loading.
Pressure drop will be a function of flow rate and pipe length.
The following equation can be used to estimate the annual
electricity costs for the fan:
Annual Fan
Electricity
Costs = 9.926 * 10~6 * (Flow Rate)(AP)(Loading Hours*2)
(July 1989 $)
where:
Flow rate = vent stream flow rate in cfm; and
AP = pressure drop in inches of water.
For the large volume of nitrogen needed to blanket the
vehicles during loading, a facility is assumed to use a nearby
air separation plant. The cost is based on the volume of
3-23
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nitrogen needed, which would equal the throughput transferred
at the facility. The following equation can be used to
estimate this cost:
Annual Cost for
N2 Blanketing = 274 * (Annual Throughput, MMgal/yr)
(July 1989$)
Indirect annual cost includes capital recovery, vehicle
vapor-tightness testing, and maintenance for the vapor
collection system. These costs are summarized in Table 3-7.
Capital recovery is calculated as:
Capital recovery = 0.163 * TCI
where the capital recovery factor (0.163) is based on an
equipment life of 10 years and an interest rate of 10 percent.
The annual cost for vehicle testing to ensure vapor tightness
is $366 per vehicle in July 1989 dollars, and the maintenance
charge for the vapor collection system is $244 per rack in
July 1989 dollars.
The costs for vapor collection and transport are combined
with the costs derived for the control device (flare or
incinerator) to calculate the overall total capital investment
and annual costs. The costs for the combustion control
devices are discussed in Section 3.1.
3.2.2 Cost Methodology for Condensers
This section discusses capital and annual costs for
packaged refrigerated surface condenser systems. These costs
were derived from vendor data.
3.2.2.1 Condenser Design Considerations Affecting Cost.
The design parameters of a refrigerated condenser system that
affect cost are the same ones that affect removal efficiency
because they determine the condensation temperature and heat
load (refrigeration tonnage) of the system. These design
parameters include the following:
• Volumetric flow rate of the VOC-containing vent
stream;
• Inlet temperature of the vent stream;
• Concentrations of the VOC's in the vent stream;
• Absolute pressure of the vent stream;
3-24
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TABLE 3-7. TOTAL ANNUAL COST FOR LOADING RACK
VAPOR COLLECTION SYSTEMS
Cost Item
Cost
(July 1989$)
Direct Annual Cost
Toll charge for N2 blanketing3 284 * (Throughput, MMgal/yr)
Electricity for fan/motor13 See equation in test
Indirect Annual Cost
Capital recovery0
Vehicle vapor-tight testingd
Maintenance for vapor
collection systemd
0.163 * TCI
$366 per vehicle
$244 per rack
aFrom Reference 21.
bFrom Reference 20.
GCapital recovery is based on an equipment life of 10 years
and an interest rate of 10 percent.
dFrom Reference 22.
3-25
-------
• Required removal efficiency of the VOC;
• Moisture content of the vent stream; and
• Properties of the VOC's in the vent stream:
heats of condensation,
heat capacities, and
vapor pressures.
The cost equations in the following sections are all
functions of condensation temperature and heat load.
3.2.2.2 Development of Capital Costs. This section
presents the procedures and data necessary for estimating
capital costs for packaged refrigerated surface condenser
systems. Total capital investment includes equipment costs
for the entire refrigerated condenser system, auxiliary
equipment costs, taxes, freight charges, instrumentation, and
installation costs. Taxes and freight charges are based on
factors described in the OCCM. Instrumentation is included
in the equipment costs for packaged systems.
Vendors of refrigerated surface condenser systems were
asked to provide cost estimates for a wide range of
applications. These vendor quotes were used to develop the
costs presented here. Table 3-8 lists equipment cost
equations for 13 condensation temperatures.
The cost estimates apply to skid-mounted systems designed
for hydrocarbon vapor recovery. The systems are operated
intermittently, allowing 30 to 60 min for defrosting by
circulation of warm brine. The achievable VOC removal
efficiencies for these systems are 90 to 95 percent.
The packaged systems include the refrigeration unit with
the necessary pumps, compressors, condensers/evaporators, and
coolant reservoirs; the VOC condenser unit and VOC recovery
tank; precooler; instrumentation and controls; and piping.
Costs for heat transfer fluids (brines) are not included. The
costs in Table 3-8 are based on a new plant installation; no
retrofit cost considerations are included because retrofit
cost factors are very site specific.
Purchased equipment costs include the packaged equipment
cost and factors for sales tax (3 percent) and freight
3-26
-------
TABLE 3-8. EQUIPMENT COST EQUATIONS FOR PACKAGED
REFRIGERATED CONDENSER SYSTEMS3
Condensation
Temperature
40
30
20
10
0
-10
-20
-30
-40
-50
-60
-70
-100
<
<
<
<
<
<
<
<
<
<
<
<
<
T
T <
T <
T <
T <
T <
T <
T <
T <
T <
T <
T <
T <
40
30
20
10
0
-10
-20
-30
-40
-50
-60
-70
Equipment Costs
(July 1989
dollars)
1,
2
2
3
5
6
3,
9
4,
5
6
8
8
11
22
813
,274
,926
,997
,016
,977
600
,450
475
,417
,824
,380
,940
,172
,248
Rb
R H
R H
R H
R H
R H
R +
R H
R +
R H
R H
R H
R H
R H
R H
+ 13,521
h 13
h 13
1- 17
\- 18
h 16
53,
h 16
51,
h 50
h 50
h 49
1- 54
h 52
!• 58
,830
,776
,465
,034
,789
834
,813
083
,111
,103
,991
,549
,141
,633
Applicable
Tonnage
R >
R >
R >
R >
R >
10 > R
R >
5 > R
0
0
0
0
0
.85
.63
.71
.44
.32
> 0.21
>
R >
R >
R >
R >
R >
R >
R >
2
1
1
1
1
0
10
0.13
5
.42
.92
.58
.25
.33
.67
aFrom References 23 and 24.
= Condenser heat load in tons of refrigeration.
3-27
-------
(5 percent). Instrumentation and controls are included in
equipment costs for the packaged units. Thus,
PEC ($) = EC (1 + 0.03 + 0.05) = 1.08 * EC
where:
PEC = purchased equipment costs; and
EC = equipment costs.
The total capital investment for packaged systems is
obtained by multiplying the purchased equipment cost by the
total installation factor25 (1.15):
TCI ($) = 1.15 * PEC
where:
TCI = total capital investment; and
PEC = purchased equipment costs.
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 installation factors. ' If an
existing condenser is removed so that a more efficient
condenser may be put in its place, the cost of the demolition
is estimated to be $284 based on 8 man-hours of labor.
3.2.2.3 Development of 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 3-9.
Direct annual costs (DAC) include labor (operating and
supervisory), maintenance (labor and materials), and
electricity. Operating labor is estimated at 1/2-hr per 8-hr
shift.2 Supervisory labor is assumed to be 15 percent of the
operating labor cost. Maintenance labor is estimated at
1/2-hr per 8-hr shift. Maintenance material costs are assumed
to equal maintenance labor costs.
Utility costs for refrigerated condenser systems include
electricity requirements for the refrigeration unit and any
pumps and blowers. The power required by the pumps and
blowers is negligible compared with the power requirements of
3-28
-------
TABLE 3-9. BASES AND FACTORS FOR ANNUAL COSTS
FOR REFRIGERATED CONDENSER SYSTEMS
Basis for Direct Annual Costs Factor
Operating labor (man-hours/8-hr shift) 0.5
Maintenance labor (man-hours/8-hr 0.5
shift)
Labor rates (July 1989 $/hr)a
Operating labor 13.20
Maintenance labor 14.50
Supervisory labor cost
(percent of operating labor cost) 15
Maintenance materials cost
(percent of maintenance labor cost) 100
Utilities (July 1989 $)a
Electricity ($/l,000 kW-hr) 50.9
Natural Gas ($/l,000 ft3) 3.03
Product recovery credit Emission Reduction *
Chemical Market Price
Basis for Indirect Annual Cost
Equipment life (years) 15
Interest rate (percent) 10
Capital recovery factor 0.1314
Taxes, insurance, administration 4
(percent of total capital investment)
Overhead (percent of total 60
labor and maintenance costs)
aBased on Reference 3.
3-29
-------
the refrigeration unit. Electricity requirements for
refrigerated condenser systems are summarized below:
Electricity (E) Temperature
1.3 kW/ton 40 °F < T
2.2 kW/ton 20 °F < T < 40 °F
4.7 kW/ton -20 °F <. T < 20 °F
5..0 kW/ton -50 °F < T < -20 °F
11.7 kW/ton -100 °F < T < -50 °F
These estimates were developed from product literature
obtained from one vendor. Electricity costs can then be
calculated from the following expression:
Ce ($/yr) = [R(tons)/0.85] * E (kW/ton) * (hr/yr) * ($/kW-hr)
where Ce = annual electricity costs;
R = refrigeration requirements;
E = electricity requirements;
and the factor 0.85 accounts for the mechanical efficiency of
the compressor.
Indirect annual costs are the sum of capital recovery
costs plus general and administrative (G&A), overhead, taxes,
and insurance costs. Overhead is assumed to equal 60 percent
of the sum of maintenance materials and operating,
supervisory, and maintenance labor.
The system capital recovery cost is based on an estimated
15-year equipment life. For a 15-year life and an interest
rate of 10 percent, the capital recovery factor is 0.1315.
Thus,
Cc ($/vr) = 0.1315 * TCI
where:
Cc = capital recovery cost; and
TCI = total capital investment.
Administrative charges, taxes, and insurance are factored
from TCI, typically 2, 1, and 1 percent, respectively.
If the condensed VOC can be directly reused or sold
without further treatment, the credits from this operation are
incorporated in the total annual cost estimates. The
following equation can be used to estimate the VOC recovery
credits:
3-30
-------
RC ($/yr) = Emission Reduction (Ib/yr) * Market Price ($/lb)
where:
RC = recovery credit ($/yr).
Data on market prices are available from sources such as the
Chemical Marketing Reporter.
Total annual costs are calculated as the sum of direct
annual costs and indirect annual costs, minus recovery
credits:
TAG = DAC + IAC - RC
where:
TAG = total annual costs;
DAC = direct annual costs;
IAC = indirect annual costs; and
RC = recovery credits.
3.2.3 Cost Methodology for Steam Stripping
This section discusses steam stripper design
considerations affecting cost and the general methodology for
developing capital and annual costs for steam strippers. A
detailed example of the application of this methodology to
cost a steam stripper for an example wastewater stream is
given in Appendix D of this volume.
3.2.3.1 Steam Stripper Design Considerations Affecting
Cost. As discussed in Chapter 2, a number of factors are
involved in the design of a steam stripper: stripper
configuration (direction of flow as well as tray vs. packed
bed design); wastewater flow rate; steam flow rate or steam-
to-feed ratio; column height and diameter; wastewater feed
temperature and pH; and the vapor-liquid partitioning of
compounds to be removed (expressed by Henry's Law constant).
This discussion is limited to sieve tray column steam
strippers with countercurrent steam flow and noncorrosive
wastewater with a typical feed temperature of 35 °C (95 °F) .
Factors affecting the steam stripper capital costs
include the column diameter, column height, and the size of
auxiliary equipment (feed tanks, feed preheater, condenser,
decanter, flame arrestor, and pumps). The column diameter and
the size of the auxiliary equipment are a function of the
3-31
-------
design wastewater feed rate. The column must be wide enough
to provide a desired pressure drop and liquid retention time
in the column using correlations developed to prevent column
flooding. As the wastewater feed rate increases, the column
diameter will increase proportionally. With increases in the
column diameter and the size of the auxiliary equipment, the
capital costs of the system increase. The height of the steam
stripper column depends upon the desired compound removal
efficiency. The column height is designed to accommodate the
height of trays necessary to achieve a desired compound
removal efficiency. Typically, higher removal efficiencies
require more trays, thereby increasing the capital cost of the
system.
The steam stripper annual costs are affected most by the
annual steam requirement, which is a function of the steam-to-
feed ratio and the wastewater feed rate. The steam-to-feed
ratio is selected to obtain the desired compound removal
efficiency. Higher compound removal efficiencies generally
require a greater steam-to-feed ratio and, therefore, result
in higher steam costs. The steam requirements for the
stripper are also a direct function of the wastewater feed
rate.
3.2.3.2 Development of Steam Stripper Capital Costs.
The total capital investment for a steam stripper system
includes purchased equipment cost, direct installation cost,
?8
and indirect installation cost. The purchased equipment
cost comprises the basic equipment cost, auxiliary piping and
equipment cost, instrumentation, freight, and sales tax. The
basic equipment cost is the sum of the price of each component
of the steam stripper system and is estimated using
engineering cost estimation techniques. ••••'• Table 3-10
presents equations for the costs of the various components of
the steam stripper system. All costs are for carbon steel
construction except for sieve trays and pumps. It was assumed
that these components would be constructed of stainless steel
because they are subject to the greatest wear and are exposed
to the harshest conditions.
3~32
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3-34
-------
The total capital investment required to install a new
steam stripper unit is calculated as a direct function of the
basic equipment cost value. The purchased equipment cost is
calculated by multiplying the basic equipment cost by an
appropriate percentage value. This percentage value and the
other multipliers discussed below are selected from ranges
recommended in cost estimation reference documents.
Piping costs are implicitly included in the direct
installation cost; however, auxiliary piping (i.e., additional
piping for the combination of wastewater streams and vapor
vent lines for storage tanks) are accounted for separately in
the purchased equipment cost. The purchased equipment cost is
used to estimate the steam stripper system direct installation
costs and indirect installation costs. Each of these costs is
calculated by multiplying the purchased equipment cost by an
appropriate percentage value. The direct installation cost
includes items such as electrical wiring, insulation,
equipment support and erection, and painting of equipment.
The indirect installation cost includes engineering,
construction and field expense, construction fee, start-up and
testing, and contingency. The sum of purchased equipment
cost, direct installation cost, and indirect installation cost
yields the total capital investment. The total capital
investment can also include costs for buildings, off-site
facilities, land, working capital, and yard improvements;
however, these costs are not typically included in the
purchased equipment cost for a steam stripper system.
Table 3-11 summarizes the equations used in estimating
total capital investment for a steam stripper system. The
total capital investment for installing a new steam stripper
system is presented in Figure 3-1 as a function of the
system's design wastewater feed rate. The total capital
investment costs for this graph were calculated using the cost
equations in Table 3-11 and are based on a steam stripper
system design which is sized for each wastewater flow rate.
The stripper system design was developed using the ASPEN
computer program. Figure 3-1 presents total capital
3-35
-------
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investment for both carbon and stainless steel construction.
Stainless steel construction costs are included for comparison
because streams with corrosive wastewater (i.e., very high or
low pH) will require a steam stripper system constructed of a
corrosion resistant material. Equipment costs for stainless
steel were developed from the same information sources as for
carbon steel. Generally, a factor for material of
construction was used for conversion of carbon steel cost to
304 stainless steel cost. As shown in Figure 3-1, the total
capital investment is a direct function of the wastewater flow
rate to the steam stripper unit, with stainless steel
construction being approximately two to seven times more
costly than carbon steel.
3.2.3.3 Development of Steam Stripper Annual Costs. The
total annual cost is the total of all costs incurred to
operate the steam stripper system throughout the year. The
annual operating costs comprise direct and indirect charges.
Direct annual costs comprise expenses incurred during normal
operation of the steam stripper process, including utilities,
labor, and maintenance activities.
Three types of utilities are used to operate the steam
stripper process unit: electricity, steam, and cooling water.
Electricity is needed to operate pumps and other electrical
components in the system. The electricity required for the
pumps is calculated using design flow rates for each pump and
assuming a developed head of approximately 37 m (120 ft) of
water and a pump efficiency of 64 percent. The steam costs
are estimated using the design steam loading of 0.096 kg
steam/£ (0.80 Ib/gal) of wastewater feed. The cooling water
cost is calculated using water requirements necessary for the
overhead primary condenser.
Other direct costs include labor and maintenance. Labor
cost was estimated using the number of hours required to
operate a steam stripper process unit (0.5 hr/shift) and a
labor rate of $13.20/hr,3 The supervisory and administrative
costs are estimated as 15 percent of operating labor. The
maintenance costs consist of labor and materials. The
-------
maintenance labor cost assumes 0.5 hr/shift operation and a
$14.50/hr labor rate. The maintenance materials cost is
assumed to equal 100 percent of maintenance labor cost.
Indirect operating expenses are incurred regardless of
the operating status of the steam stripper system. The cost
of overhead is estimated to be 60 percent of all labor and
maintenance costs. The remaining components of the indirect
annual costs are taken as percentages of the total capital
investment. The capital recovery for the steam stripper
system is based on a 15-year equipment life at an interest
rate of 10 percent.
Another element of the total annual cost that is included
as a benefit in this estimate is the credit for the recovered
chemicals. This factor accounts for any cost credits which
result from the organic chemicals being recovered from the
overheads stream. There are several alternatives for handling
the recovered organic chemicals. If steam is produced
on-site, the recovered organic chemicals can be used as fuel
for the existing boiler. The money saved by not having to
purchase conventional fuels (i.e., fuel oil or natural gas) is
the recovery credit. In this situation, the value of the
recovered chemicals is equal to the fuel value only. Another
option is to reuse the recovered chemicals in the
manufacturing process. In some cases the organic chemicals
can be recycled directly to the process; in other cases the
organic chemicals must be separated by distillation before
reuse. The costs saved in reducing the purchase of raw
materials is the recovery credit and is valued at the cost of
the chemicals; however, this cost savings may be offset by the
cost of distilling the recovered organic chemicals. Another
option is to sell the recovered organics to a chemical
manufacturer who will recover the individual components in the
waste organic stream. However, in cases where a cost-
effective use for the recovered organics does not exist, the
plant will have to pay for disposal of the collected organic
chemicals. There will be no cost savings in this case; in
fact, an additional cost for disposal may be incurred.
3-39
-------
This cost estimation methodology assumes that the
recovered organic chemicals are used as fuel for an existing
boiler rather than recycled to the process or sold to a
chemical manufacturer. To calculate the recovery credit, the
heat content of the recovered chemicals was estimated based on
their composition.
Table 3-12 summarizes the equations discussed above which
are used to estimate total annual cost for a steam stripper
system. The annual unit operating cost ($/£) for the steam
stripper is calculated by dividing the total annual cost
($/Yr) by tne annual wastewater throughput (£/yr). The annual
unit operating cost for installing and operating a steam
stripper system is compared to the system's design wastewater
feed rate in Figure 3-2. These annual unit operating costs
were calculated using the cost equations in Table 3-11 and are
based on a stripper system design which is sized for each
wastewater flow rate. The stripper system design was
developed using the ASPEN computer program. The costs for
both carbon and stainless steel construction are presented.
As shown in Figure 3-2, the annual unit operating cost is an
indirect function of the wastewater feed rate to the steam
stripper unit, with stainless steel construction being
approximately 1.5 times more costly than carbon steel.
3.3 COST METHODOLOGY FOR STORAGE TANK IMPROVEMENTS
3.3.1 Design Considerations Affecting Cost
As described in Chapter 2, there are many parameters
affecting the emission control efficiency of internal floating
roof tanks. Some of the parameters are chemical specific
while others are related to the floating deck. The deck
design considerations that affect cost are:
• The number and type of rim seals;
• The type of deck fittings (i.e., controlled or
uncontrolled); and
• The type of deck seams (i.e., bolted or welded).
For this analysis, when a new floating roof is installed
on a fixed roof tank, it is assumed that the deck is bolted
and has a liquid-mounted primary seal and controlled fittings.
3-40
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Typical existing internal floating roofs have vapor-mounted
primary seals and uncontrolled deck fittings. Some tanks also
have secondary seals. These existing internal floating roof
tanks are assumed to be upgraded by the addition of a
secondary seal (when necessary) and controlled deck fittings.
The costs presented in this section were taken from the
EPA report "Control of Volatile Organic Compound Emissions
from Volatile Organic Liquid Storage in Floating and Fixed
Roof Tanks,"3 which contains costs and cost equations for
several different deck designs. For consistency with other
sections in this chapter, these costs have been updated to
July 1989 dollars.
3.3.2 Development of Capital Costs
3.3.2.1 Installing a Floating Roof. Before an internal
floating roof can be installed, the tank must be emptied,
cleaned, and degassed. During degassing, the tank is emptied
of all VOC vapors by replacing the VOC-laden air in the tank
with fresh air. Fans or blowers are used to draw in the fresh
air and exhaust the VOC vapors. The cost for this activity in
terms of tank size is as follows:
Cleaning/
Degassing (July 1989 $) = 7.61 * (Tank Size [gal])°-5132
of tank
The cost of installing a floating roof depends upon the
type of deck and seal system selected. The following equation
estimates the cost of installing a new bolted, floating deck
having a liquid-mounted primary seal and controlled deck
fittings:
New Floating
Roof = 509 * (Tank Diameter [ft]) + 1160
(July 1989 $)
This equation includes the retrofit cost of cutting vents or
openings necessary for modifying the tank.
In some cases, an existing refrigerated condenser system
may be removed before an internal floating roof is installed.
The cost of this activity is estimated to be $284 per tank
based on 8 man-hours of labor.27
3-43
-------
3.3.2.2 Upgrading an Existing Internal Floating Roof
Tank. Typical existing internal floating roof tanks have a
bolted deck, vapor-mounted primary seal and uncontrolled deck
fittings. To reduce emissions from such tanks, the deck may
be upgraded by adding a secondary seal and controlled deck
fittings (see Table 2-2).
The cost of adding a secondary seal to an existing
floating roof is given in terms of tank diameter:
Addition of
a Secondary (July 1989 $) = 95.1 * (Tank Diameter [ft])
Seal
The cost of controlled fittings is about 4 percent of the cost
•70
of a new deck having a vapor-mounted primary seal. In terms
of tank diameter, this cost is calculated as follows:
Addition of
Controlled
Deck (July 1989 $) = 16 * (Tank Diameter [ft]) + 46
Fittings
For this analysis, it is assumed that existing internal
floating roof tanks will not be upgraded until a scheduled
cleaning and degassing. Thus, the cleaning/degassing cost
will not be a part of the control costs required for upgrading
existing internal floating roof tanks.
3.3.3 Development of Total Annual Costs
The annual cost without product recovery credit is the
sum of annualized capital cost, operating costs, and costs for
taxes, insurance, and administration. Assuming an equipment
life of 10 years and an interest rate of 10 percent, the
capital recovery factor equals 0.163. Operating costs include
the yearly maintenance charge of 5 percent of the capital cost
and an inspection charge of 1 percent of the capital cost.
Taxes, insurance, and administration are assumed to equal
39
4 percent of the capital cost.
The total annual cost, TAG, also accounts for the value
of any recovered product. Product recovery credit is
calculated by multiplying the market value of the chemical by
the emission reduction achieved by the tank improvements.
3-44
-------
The equation below summarizes the calculation of total
annual cost. The first term represents the capital recovery
costs, operating costs, and costs for taxes, insurance, and
administration and the second term represents product recovery
credit:
Capital Emission Market
TAG (July 1989 $) = (Costs * 0.263)-(Reduction * Price )
where:
TAG = total annual cost.
3.4 COST METHODOLOGY FOR EQUIPMENT LEAK CONTROL TECHNOLOGIES
Emissions from different equipment types discussed in
Chapter 2 can be reduced by either control equipment or work
practices. Costs for reducing emissions from compressors,
open-ended lines, sampling connections, and pressure relief
devices are developed assuming control equipment will be
installed. Costs for reducing emissions from pumps, valves,
and connectors are developed assuming emissions will be
reduced through work practices. The general methodology is
developed to determine base costs for the control of emissions
for each type of equipment. These base costs per component
are then used with plant equipment counts to develop capital
and annual costs for the control of emissions from equipment
leaks for each plant.
3.4.1 Control Equipment
This section presents the costs associated with the
purchase and installation of specified equipment for the
control of VOC emissions from compressor seals, open-ended
lines, sampling connections, and pressure relief devices. The
costs for controlling emissions from product accumulator
vessels with a combustion device are the same as the costs for
controlling emissions from process vents, which are described
in Section 3.1.
The base costs for all of these devices were developed
assuming the same equipment requirements as stated for these
devices in the Additional Information Document for Fugitive
Emission Sources of Organic Compounds (SOCMI Fugitives AID).
3-45
-------
Base cost information presented in the Background Information
Document for Benzene Fugitive Emissions (Benzene BID) was also
41
reviewed. The base cost estimates included here were made
using Richardson Process Plant Construction Estimation
42
Standards where possible and by contacting vendors for
current prices. It was not possible to use these estimation
techniques for all control equipment, however. The prices for
control equipment for which estimates or quotations were not
available were updated from the SOCMI Fugitives AID. All
costs were converted to July 1989 dollars using the ratio of
the CE Plant Cost Indices. These base cost estimates are
presented in Table 3-13.
3.4.1.1 Compressors. The equipment for which cost
estimates were developed for compressor seals includes a heavy
liquid or non-VOC barrier fluid with a degassing reservoir
which is connected by a closed vent system to an additional
control device (e.g., flare) or vapor recovery header.4
The base cost is calculated for the purchase of 122 m of
2-in. schedule 40 steel pipe, three 2-in. steel plug valves,
40
and one metal gauze flame arrestor for each compressor. The
costs for piping and valves are estimated to be $1,090 and
44
$2,440, respectively. The cost for a metal gauze flame
arrestor is estimated to be $862. A total of 82 hr of
maintenance labor would be required for installation. The
labor cost is estimated to be $22.50/hr ,46'47 The total
installed base cost for a closed vent system is $6,242 per
compressor.
These costs are based on connecting the closed vent
system to an existing enclosed combustion device or vapor
recovery header and do not reflect the cost of adding a
control device specifically to control the degassing vents.
The actual base cost will vary depending on the length of pipe
and number of valves required for each compressor within each
plant.
3.4.1.2 Open-Ended Lines. A cap, plug, blind flange, or
second valve will be installed on all open-ended lines to
prevent emissions through the open end. The base cost is
3-46
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TABLE 3-13.
BASE COST DATA FOR EQUIPMENT LEAK
CONTROL DEVICES3
Equipment Type: Item
Cost (July
1989 $)
Basis
Compressors: closed vent
system for degassing
reservoirs
• 122 m of 5.1-cm 1,090
schedule 40 steel
pipe
• Installation (66 hr) 1,490
• Three 5.1-cm cast 2,440
steel plug valves
• Installation (6 hr) 135
• One metal gauze flame 862
arrestor
• Installation (10 hr) 225
TOTAL: 6,242
Open-ended lines: 70
One 2.5-cm gate valve
i?
• Installation (1.4 hr)
TOTAL: 102
Sample connections:
closed purge system
• 6 m of 2.5-cm 25
schedule 40 steel
pipe
• Installation (1 hr) 23
• Three 2.5-cm ball 267
valves, carbon steel
• Installation (6 hr) 95
TOTAL: 409
$261.03/100 ft pipe
CE Index(356.0/342.5)
16.6 hr/100 ft pipe
Labor $22.50/hr
$783.75/valve
CE Index(356.0/342.5)
2 hr/valve
Labor $22.50/hr
10/90 quote of
$869.00
CE Index(356.0/358.7)
Labor $22.50/hr
$67.50/valve
CE Index(356.0/342.5)
1.4 hr/valve
Labor $22.50/hr
$120.64/100 ft pipe
CE Index(356.0/342.5)
12.6 hr/100 ft
Labor $22.50/hr
$85.60/valve
CE Index(356.0/342.5)
1.4 hr/valve
Labor $22.50/hr
3-47
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TABLE 3-13. BASE COST DATA FOR EQUIPMENT
LEAK CONTROL DEVICES3 (CONCLUDED)
Equipment Type; Item
Cost
(July
1989 $)
Basis
Pressure relief seals
(new unit): rupture
disk assembly
• One 7.6-cm 78
stainless steel
rupture disk
• One 7.6-cm carbon 579
steel rupture disk
holder
• One 0.6-cm 24
pressure gauge,
dial face
• One 0.6-cm bleed 40
valve, carbon
steel
• Installation 360
(16 hr)
• One 7.6-cm gate 401
valve
• Installation 34
(1.5 hr)
• Offset: one 10.2- 49
cm tee and one
10.2-cm elbow
• Installation 191
(8.5 hr)
TOTAL: 1,755
Pressure relief seals
(retrofitted unit): 1,755
new assembly plus a
new relief valve.
• One 7.6-cm 1,950
stainless steel
body and trim
relief valve
• Installation 225
(10 hr)
TOTAL: 3,930
10/90 quote of $235/3-pak
CE Index (356.0/358.7)
10/90 quote of $583/holder
CE Index (356.0/358.7)
Update from SOCMI NSPS AID
CE Index (356.0/224.7)
Update from SOCMI NSPS AID
CE Index (356.0/224.7)
Labor $22.50/hr
$385.60/valve
CE Index (356.0/342.5)
1.5 hr/valve
Labor $22.50/hr
$35.63/tee, $11.73/elbow
CE Index (356.0/342.5)
5.7 hr/tee,
2.8 hr/elbow
Labor $22.50/hr
Update from SOCMI NSPS AID
CE Index (356.0/224.7)
Labor $22.50/hr
aCosts and hours of installation from Reference 46.
3-48
-------
estimated based on the purchase price of a 1-in. gate valve of
$70 and 1.4 hr of labor for installation for each open-ended
line. The installed base cost for each open-ended line is
estimated to be $102, including labor.
The actual base cost for each open-ended line will vary
based on the type of closure chosen. Caps, plugs, and blind
flanges are all expected to cost less than a second valve for
the same size line but cannot be used in all cases. For
example, blind flanges can only be installed on lines larger
than 1 in. in diameter.
3.4.1.3 Sampling Connection Systems. Sampling
connection systems should be equipped with closed-loop purge
systems to prevent losses of process fluid to the atmosphere.
The base cost for each closed-loop purge system is based on
the installation of 6 m of 1-in. schedule 40 steel pipe and
40
three 1-in. carbon steel ball valves. Base cost estimates of
44
the piping and valves are $25 and $267, respectively. The
estimated total base cost after installation is $409 per
sample connection.
3.4.1.4 Pressure Relief Devices. The emissions from
pressure relief valves in gas/vapor service can be controlled
by installing a rupture disk in the line upstream from the
relief valve. This requires the purchase and installation of
a rupture disk and a rupture disk holder for each pressure
relief valve. The purchase cost of these components are
/ 0
estimated at $78 and $579, respectively.
Performing maintenance on the rupture disk assembly
during plant operation requires a valve upstream from the
rupture disk, a pressure gauge, and a bleed valve.40 The valve
is estimated to cost $401.44 The pressure gauge and bleed
valve are estimated to cost $24, and $40, respectively.40 To
protect the pressure relief valve from damage caused by
fragments of the rupturing disk at the time of pressure
relief, a piping offset is typically installed.40 This offset
comprises one 3-in. tee costing approximately $36 and one
40
3-in. elbow costing approximately $12. Installing these
components is estimated to require 26 hr, for a total
3-49
-------
installed cost estimate of $1,755 per assembly. This
estimated base cost per pressure relief device is for new
plants.
Existing plants which are required to retrofit pressure
relief valves with this new equipment also typically require a
new pressure relief valve. A 3-in. stainless-steel body and
trim relief valve costing $1,950 and requiring 10 hours to
40
install is assumed. This increases the total estimated base
cost for existing plants to $3,930 per pressure relief device.
3.4.2 Leak Detection and Repair Techniques
This section presents the costs associated with the
initiation and continuation of a leak detection and repair
program for the control of VOC emissions from pumps, valves,
and connectors. Because leak frequency data for agitators are
insufficient, no estimates of repair costs were made for this
analysis. The base costs for this program are discussed in
two parts: the leak detection program and the repair program.
3.4.2.1 Leak DetectionJProgram. The various methods
used to identify equipment leaks are discussed in Section 2.4.
This section discusses the method used to develop base costs
for a leak detection program based on equipment counts and
monitoring frequencies. In the development of these base
costs, only the methods of an individual component survey by
instrument inspection and by visual inspection are considered.
The procedure used to estimate the number of leaks detected
annually based on equipment counts, monitoring frequencies,
and leak frequencies is also discussed in this section. "Leak
frequency" is the fraction of all pieces of equipment of each
type which have screening values above a given action level,
measured at any given time. Leak frequency must be considered
in the context of this action level. For example, leak
frequencies provided in the SOCMI Fugitives AID were based on
an action level of 10,000 ppmv.
For the purposes of cost estimation, vendors who perform
leak detection monitoring by instrument inspection were
contacted.49'50 Based on the information provided by these
contractors, the average cost of the first individual
3-50
-------
component survey was determined to be $2.50 per component
monitored. An average cost of $2.00 per component monitored
per survey was determined for all surveys after the initial
survey. The cost of the initial survey is higher because of
difficulties in initially locating all components, increased
numbers of leaking components, and increased time needed to
number and tag equipment. ' The cost of an initial
individual component survey will be different for each process
unit. The cost will depend on the number of equipment
components to be monitored and the actual initial leak
frequencies of the process unit.
The monitoring costs cover the equipment rental and labor
charges for a contractor's survey team to "screen" each piece
of equipment using a portable monitoring instrument to
determine which pieces of equipment require maintenance. This
monitoring team would also tag leaking components and make a
first attempt to repair the leak by tightening the packing
gland on valves or by some other technique which does not
require removing the equipment from service. If a process
unit chose to conduct monitoring internally, it is assumed the
cost would not be greater than the contracted cost.
For the purpose of cost estimates, different monitoring
frequencies for each type of equipment being surveyed are
assumed. The monitoring frequency assumed is monthly for
pumps, agitators, and valves, and annually for connectors.
These frequencies may vary from plant to plant, based on the
number of pieces of equipment that have been identified for
maintenance at each plant. For example, if a plant has under
2 percent of its valves requiring maintenance, that plant may
43
monitor valves quarterly.
It is assumed for this cost analysis that each process
unit is required to purchase one portable monitoring
instrument. Very large process units may require more than
one instrument, but plants with multiple process units may not
need to purchase one for each process unit. This instrument
would serve as a backup to the contractor's instrument and
would be used to verify that components identified as
3-51
-------
requiring repair by the contractor have been repaired. The
cost of each monitoring instrument adjusted to July 1989
dollars is estimated at $6,500.51
3.4.2.2 Repair Program. This section discusses the
method used to develop base costs for the repair of equipment
determined to be leaking by a monitoring team but not
successfully repaired during initial repair efforts. These
leaking components are required to be repaired within 15 days
of leak detection. The costs for leak repair are based on the
40
leak repair costs estimated in the SOCMI Fugitives AID.
These base costs can be used with average leak frequencies and
equipment counts to determine an annual operating cost for
leak repairs.
The average number of leaks detected by the leak
detection program can be estimated based on the number of
components monitored, the monitoring frequency, and the leak
frequency. An estimate of the average number of leaks
detected is necessary to estimate the cost of a repair
program. By multiplying the leak frequency for each type of
equipment by the number of components, the estimated number of
leaks can be determined. Leak frequencies are estimated for
the initial leak detection survey and for all subsequent
surveys.
Initial leak frequencies for valves and connectors, based
on the number of pieces of equipment with screening values
over 10,000 ppmv, were estimated from data in the EPA 24-Unit
Study.52 Initial leak frequencies for gas valves, light
liquid valves, and connectors are 11.4 percent, 6.5 percent
and 2.1 percent, respectively. The initial leak frequency for
pumps is assumed to be two times the base performance level
allowed under Phase III of the negotiated regulation. The
leak frequency for pumps from the 24-Unit Study was not used
in this analysis because the value (based on 10,000 ppmv) was
below the final performance level for pumps provided in the
negotiated regulation. Using the leak frequency from the
24-Unit Study would underestimate the number of pump seals to
be repaired, because the action level under the negotiated
3-52
-------
regulation is reduced to 1,000 ppmv. The frequencies for all
surveys subsequent to the initial survey are estimated to be
the base performance levels allowed under the final phase of
the negotiated regulation, which are 2 percent for all valves,
43
0.5 percent for connectors, and 10 percent for pumps.
All of the costs and assumptions discussed above for the
leak detection program are summarized in Table 3-14. These
values can be substituted into the equations presented in
Table 3-15 to determine the cost of the leak detection
program. Repair of each equipment type is described below.
3.4.2.2.1 Pumps. The SOCMI Fugitives AID presents costs
for leak detection and repair of pump seals. The LDAR model
was used to estimate the number of pump seals with screening
values above 10,000 ppmv, which would therefore require
repair. Because repair of pump seals demands more extensive
maintenance efforts than possible by a monitoring team,
follow-up repair was assumed for 100 percent of the pump seals
predicted to have screening values over 10,000 ppmv. This
assumption is not consistent with the analysis needed to
estimate cost impacts of the negotiated regulation.
According to the negotiated regulation, the leak
43
definition for pump seals will be 1,000 ppmv of VOC.
However, the negotiated regulation also states that only pump
seals with leakage through the seal with a monitored
concentration of greater than or equal to 2,000 ppmv of VOC
will require repair. Therefore, this analysis of cost impacts
considers the number of pump seals with leakage equivalent to
a monitored concentration above 2,000 ppmv.
To estimate the actual percentage of leaking pump seals
that will require repair, screening value data from
19 ethylene oxide and butadiene plants were evaluated. By
counting the number of leaking pumps that had screening values
greater than 2,000 ppmv VOC, it was determined that
approximately 75 percent of all leaking pumps would require
subsequent repairs.
It is assumed that all pumps requiring repair will
40
require a new seal. The cost for each replacement single
3-53
-------
TABLE 3-14. BASE COSTS AND ASSUMPTIONS FOR A
LEAK DETECTION PROGRAM
Subcontractor monitoring fee:
Initial monitoring
All subsequent monitoring
$2.50/component
$2.00/component
Monitoring frequencies:
Pumps
Connectors
Valves
Monthly
Annually
Monthly
Initial leak frequencies:
Pumps
Connectors
Gas/vapor valves
Light liquid valves
20 percent
2.1 percent
11.4 percent
6.5 percent
Leak frequencies (subsequent to initial
monitoring):
Pumps
Connectors
Gas/vapor valves
Light liquid valves
10 percent
0.5 percent
2 percent
2 percent
Visual inspection of pumps:
Time
Labor charge
30 sec/pump
$22.50/hr
Organic vapor detection instrument:
$6,500
3-54
-------
TABLE 3-15. EQUATIONS FOR DETERMINING COSTS AND NUMBER OF
LEAKS FOR A LEAK DETECTION PROGRAM
Initial Monitoring Costs:
Cost of _ /Number of Pieces\ + / Subcontractor \
Monitoring \ of Equipment / \ Monitoring Fee/
Subsequent Monitoring Costs:
Cost
/Number of\ /Subcontractor^
= Pieces of * Monitoring *
Monitoring Equipment) ( Fee )
Monitoring
Frequency
(x/yr)
Initial Monitoring Number of Leaks:
Initial = /Number of Pieces\ ^ / Leak \ [ Monitoring 1
Leaks \ of Equipment ) ^Frequency/ [Frequency (x/yr)j
3-55
-------
seal is estimated to be $180, which includes a 50 percent
credit for the old seal.40 The cost of a single seal is used
for estimating replacement seal costs for all pumps. Double
mechanical seals are not used because they are more expensive
and they could be exempted from routine leak detection and
repair by using a barrier fluid system with degassing
reservoir connected to a closed vent system. The time
required for these repairs is estimated to be 16 hr for each
tie 1
46,47
40
pump. The labor costs for this repair time is estimated at
$22.50/hr.
3.4.2.2.2 Valves. In the SOCMI Fugitives AID, it was
assumed that 25 percent of all valves identified as leaking
would require repairs beyond any initial efforts made by the
monitoring team. It was assumed that the remaining 75 percent
of the leaking valves had been repaired by the initial efforts
40
of the monitoring team. This assumption has been retained
for this analysis.
The time of the additional repairs is estimated to be
40
4 hours for each valve. The labor costs for this repair time
is estimated at $22 .50/hr.46'47 Costs for replacement valve
seals are not included here. Replacement seals are considered
to be covered by routine plant maintenance because these costs
are typically low and not all valves requiring repair will
require replacement seals.
3.4.2.2.3 Connectors. In the SOCMI Fugitives AID, it
was assumed that 25 percent of all connectors identified as
leaking would require repairs beyond any initial efforts made
by the monitoring team. It was assumed that the remaining
75 percent of the leaking connectors had been repaired by the
initial efforts of the repair team. Although the leak
definitions have changed, this assumption has been retained
for this analysis.
The time required for additional repairs is estimated to
40
be 2 hours for each connector. The labor cost for this
repair time is estimated at $22.50/hr.46'4 Costs for
replacement connector seals are not included here.
Replacement seals are considered to be covered by routine
3-56
-------
plant maintenance because these costs are typically low and
not all connectors requiring repair will require replacement
seals.
3.4.3 Capital Costs
The capital costs for the control of emissions from
equipment leaks include capital expenditures for all control
equipment and the purchase cost of a portable organic vapor
detection instrument. The cost of the initial leak detection
and repair program is also included as a capital expense
because the higher costs of the initial monitoring are only
incurred in the first year.
The capital costs for control equipment are estimated by
multiplying the base cost for each type of equipment
(Table 3-13) by the equipment count for that type of equipment
in the process unit. The purchase cost of a portable organic
vapor detection instrument is $6,500. The initial LDAR
program costs are developed as discussed in Section 3.4.2 of
this volume. An additional 40 percent is also added to the
initial leak detection and repair program costs to cover
40
administration and support expenses.
3.4.4 Annual Costs
The annual costs for the control of VOC emissions from
equipment leaks have been divided into five sections:
annualized capital costs, annual maintenance charges, annual
miscellaneous charges, costs for leak detection and repair,
and recovery credits. Each of these sections is discussed in
detail below.
3.4.4.1 Annualized Capital Costs. The annualized
capital costs are calculated by taking the appropriate factor
from Table 3-16 and applying it to the corresponding capital
cost. The capital recovery factors are calculated using the
equation:
CRF =
(1 + i)n - 1
where:
CRF = capital recovery factors;
3-57
-------
TABLE 3-16.
DERIVATION OF ANNUAL COSTS FOR CONTROL
OF EQUIPMENT LEAKS^
Capital recovery factor for
capital charges
Rupture disks
Monitoring instruments
Other control equipment
Initial labor
Initial replacement seals
Annual maintenance charges
• Control equipment
• Replacement pump seals
• Monitoring instrument
Annual miscellaneous charges
(taxes, insurance,
administration)
• Control equipment
• Monitoring instruments
• Replacement pump seals
Annual operating costs
• Contractor monitoring fee
• Subsequent repair labor
charges
• Administrative and support
Recovery credits
• VOC raw material/product cost
• HAP raw material/product cost
0.58 * Capital
0.23 * Capital
0.163 * Capital
0.163 * Costs
0.58 * Costs
0.05 * Capital
Estimated Number of Leaks per
Process Unit * 0.75 *
$180/Seal
$4,280
0.04 * Capital
0.04 * Capital
0.80 * Annual Maintenance
Charge
$2.00/Component/Monitoring
22.50/Hr
1.4 * (Contractor Fee +
Subsequent Repair Charges)
Market Valueb * Mass of VOC
Saved
Market Valueb * Mass of HAP
Saved
aFrom Reference 46.
^Depends on the specific chemicals used and the composition of
the process stream.
3-58
-------
i = interest rate, expressed as a decimal; and
n = economic life of the component, in years.
The interest rate is 10 percent. The expected life of the
monitoring instrument is 6 years, compared to 10 years for
control equipment. Rupture disks and pump seals are assumed
40
to have a 2-year life.
3.4.4.2 Annual Maintenance Charges. The annual
maintenance charge for control equipment is calculated by
multiplying the appropriate cost by 5 percent.46 This cost
includes the expenses for labor, materials, and supervision to
40
keep the control equipment in efficient operating condition.
The annual maintenance charge for the portable monitoring
instrument is $4,280/yr. This cost was updated from the
SOCMI Fugitives AID using the CE Plant Cost Index ratio.
The cost of replacement pump seals for the LDAR program
is considered to be a maintenance expense. It is calculated by
multiplying the replacement seal cost of $180 per seal by the
number of pump leaks repaired per year.
3.4.4.3 Annual Miscellaneous Charges. The annual
miscellaneous charges cover taxes, insurance, administration,
and other fees associated with operations. The miscellaneous
charges for control equipment and for the portable organic
vapor detection instrument are calculated by applying the
factor of 4 percent from Table 3-16 to the appropriate capital
costs. 6 The miscellaneous charge for replacement pump seals
is calculated as 80 percent of the annual maintenance charge
for pump seals. The calculation is based on the miscellaneous
charge for control equipment being 4 percent of the capital
cost and the annual maintenance charge being 5 percent of the
capital cost. The miscellaneous charge is 80 percent of the
annual maintenance charge.
3.4.4.4 Leak Detection and Repair Operating Costs. The
annual operating costs for the LDAR program are developed as
discussed in Section 3.4.2. The cost of administration and
support of the LDAR program is estimated as an additional
40 percent of the leak detection and repair costs. This
3-59
-------
factor is included in Table 3-16, which summarizes the annual
LDAR operating costs.
3.4.4.5 Recovery Credits. The recovery credit is the
value of the VOC which would have been lost through emissions
if equipment leaks were not controlled. This represents a
savings to the process unit because recovery will reduce the
amount of raw material to be purchased or will increase the
amount of product to be sold, at no additional expense to the
process unit.
The recovery credit is calculated by multiplying the raw
material/product cost by the amount of the reduction in VOC
emissions from the process unit. This raw material/product
cost and amount of pollutant emissions decrease will be
different for each process unit and must be determined
specifically for each process unit.
3-60
-------
3.5 REFERENCES
1. U.S. Environmental Protection Agency, Office of Air
Quality Planning and Standards. OAQPS Control Cost
Manual. Fourth Edition. EPA-450/3-90-006. Research
Triangle Park, NC. January 1990.
2. Code of Federal Regulations, Title 40 Part 60.18.
General Control Device Requirements. Washington, DC.
U.S. Government Printing Office. January 21, 1986.
3. Memorandum from Zukor, C., Radian Corporation, to HON
project file. August 23, 1991. Documentation of cost
factors used by the HON project.
4. U.S. Environmental Protection Agency, Air and Energy
Engineering Research Laboratory. Handbook—Control
Technologies for Hazardous Air Pollutants.
EPA-625/6-86-014. Research Triangle Park, NC.
September 1986.
5. Hall, R.S., M.W. Vatavuk, J. Matley. Estimating Process
Equipment Costs. Chem. Eng. November 21, 1988. p. 66-
75
6. Vatavuk, W. Pricing Equipment for Air Pollution Control.
Chem. Eng. May 1990. p. 126-130.
7. Telecon. Stone, O.K., Radian Corporation, with Dowd, E.
ARI Technology. January 18, 1990. Incinerator sizes and
turndown.
8. Blackburn, J.W. (IT Enviroscience). Control Device
Evaluation: Thermal Oxidation. In: Organic Chemical
Manufacturing, Volume 4: Combustion Control Devices.
Report 1. U.S. Environmental Protection Agency, Research
Triangle Park, NC. Publication No. EPA-450/3-80-026.
December 1980. pp. III-l to III-5.
9. Memorandum from Scott, K., Radian Corporation, to HON
project file. February 5, 1992. Composition of design
molecule.
10. Memorandum and attachments from Farmer, J.R., EPA/ESD, to
Ajax, B. et al. August 22, 1980. Thermal incinerators
and flares.
11. Danielson, John, A. Air Pollution Engineering Manual,
Second Edition. Air Pollution Control District County of
LA. U.S. Environmental Protection Agency, Office of Air
Quality Planning and Standards. Research Triangle Park,
NC. May 1973.
3-61
-------
12. Treybal, R.E. Mass Transfer Operations, Third Edition.
New York, McGraw-Hill Book Company.
13. Memorandum from Barbour, W., Radian Corporation, to HON
project file. April 20, 1990. Estimating liquid and
vapor schmidt numbers for acid streams.
14. Memorandum from Ferrero, B., Radian Corporation to HON
project file. February 5, 1992. Estimating liquid to
vapor flow rate ratios in scrubber columns.
15. Memorandum from Ferrero, B., Radian Corporation, to HON
project file. February 5, 1992. Development of the
slope of the equilibrium line and the absorption factor
for acid gas scrubber design.
16. Memorandum. Probert, J.A., Radian Corporation, to HON
project file. September 4, 1991. Cost equations for fan
and motor in vapor collection system for transfer loading
racks.
17. U.S. Environmental Protection Agency, Office of Air
Quality Planning and Standards. Bulk Gasoline Terminals
—Background Information for Promulgated Standards.
EPA-450/3-80-038b. Research Triangle Park, NC. August
1983. p. B-5.
18. Telecon. Ocamb, D.D., Radian Corporation, with Barbe,
B. , Exxon Corporation, Baton Rouge, LA. April 13, 1989.
Cost for converting top loading railcars to bottom
loading.
19. Memorandum from Olsen, T.R., Radian Corporation, to HON
project file. August 21, 1991. Loading techniques
utilized in the SOCMI.
20. Memorandum from Probert, J.A., Radian Corporation, to HON
project file. August 28, 1991. Annual electricity costs
for motors that power fans in a vapor collection system
for transfer loading racks.
21. Memorandum from Probert, J.A., Radian Corporation, to HON
project file. August 28, 1991. Costs for installing a
nitrogen blanketing system on a transfer rack.
22. Memorandum from Probert, J.A., Radian Corporation, to HON
project file. September 3, 1991. Costs for installing a
vapor collection system on a transfer rack.
23. Edwards Engineering Corporation. Equipment
Specifications for Solvent Vapor Recovery Units, Low
Temperature Packaged Liquid Chillers, and Vapor
Condensers. New York. September 1990.
3-62
-------
24. U.S. Environmental Protection Agency, Office of Air
Quality Planning and Standards. OAQPS Control Cost
Manual. Fourth Edition. Supplement 1. Chapter 8.
EPA-450/3-90-006a. Research Triangle Park, NC.
January 1992.
25. Letter and attachments from Waldrop, R., Edwards
Engineering Corp., to Barbour, W., Radian Corporation.
September 28, 1990. Response to letter of inquiry
regarding cost of refrigerator surface condenser systems.
26. Vatavuk, W.M., and R.B. Neveril. Estimating Costs for
Air Pollution Control Systems, Part II: Factors for
Estimating Capital and Operating Costs. Chem. Eng.
November 3, 1980. pp. 157-162.
27. Memorandum from Probert, J.A., Radian Corporation, to
Docket No. A-90-21. September 10, 1990. Costs for
condenser removal.
28. U.S. Environmental Protection Agency, Office of Air
Quality Planning and Standards. EAB Control Cost Manual
Chapter 2: Cost Estimating Methodology. 4th Edition.
Draft. Research Triangle Park, NC. March 1989. p. 2-5
to 2-8.
29. Corripio, A.B., K.S. Chrien, and L.B. Evans. Estimate
Costs of Heat Exchangers and Storage Tanks via
Correlations. Chem. Eng. January 25, 1982. p. 145.
30. Peters, M.S., and K.D. Timmerhaus. Plant Design and
Economics for Chemical Engineers. Third Ed. New York,
McGraw-Hill Book Company. 1980. pp. 768-773.
31. Mulet, A., A.B. Corripio, and L.B. Evans. Estimate Costs
of Distillation and Absorption Towers via Correlations.
Chem. Eng. December 28, 1981. p. 180.
32. Ref. 30, p. 572, Figure 13-58.
33. Telecon. Gitelman, A., Research Triangle Institute, with
Oakes, D., Hoyt Corporation. September 9, 1986. Cost of
flame arresters.
34. Richardson Engineering Services, Inc. Process Plant
Construction Estimating Standards, Volume 3.
Section 15-40. Mesa, AZ. 1988.
35. Memorandum from Peterson, P., Research Triangle
Institute, to S. Thorneloe, EPA/CPB. January 18, 1988.
Basis for steam stripping organic removal efficiency and
cost estimates used for source assessment model (SAM)
analysis.
36. Ref. 28, p. 4-27.
3-63
-------
37. U.S. Environmental Protection Agency. Control of
Volatile Organic Compound Emissions from Volatile Organic
Liquid Storage in Floating and Fixed Roof Tanks. Draft
Report. Office of Air Quality Planning and Standards.
Research Triangle Park, NC. June 1984.
38. U.S. Environmental Protection Agency, Office of Air
Quality Planning and Standards. VOC Emissions from
Volatile Organic Liquid Storage Tanks—Background
Information for Promulgated Standards.
EPA-450/3-81-003b. Research Triangle Park, NC.
January 1987.
39. Ref. 37, p. 5-6.
40. U. S. Environmental Protection Agency. Fugitive
Emissions of Organic Compounds—Additional Information on
Emissions, Emission Reductions, and Costs.
EPA-450/3-82-010. Research Triangle Park, NC.
April, 1982. Section 5.
41. U. S. Environmental Protection Agency, Office of Air
Quality Planning and Standards. Benzene Fugitive
Emissions—Background Information for Promulgated
Standards. EPA-450/3-80-032b. Research Triangle Park,
NC. June 1982.
42. Ref. 34, Sections 15-42, 15-43, 15-55.
43. National Emission Standards for Hazardous Air Pollutants;
Announcement of Negotiated Regulation for Equipment
Leaks. Federal Register, Vol. 56, No. 44, pp. 9315-9339.
Washington, DC. Office of the Federal Register.
March 6, 1991.
44. Ref. 34, Section 15-42.
45. Letter from Caracciolo, D. , NAO Inc.', to Whitt, D. ,
Radian Corporation. October 31, 1990. Response to
request for cost of NAO flame arrestor.
46. Memorandum from Whitt, D., and K. Hausle, Radian
Corporation, to Markwordt, D., EPA/CPB. February 21,
1992. Final cost impacts analysis for HON equipment
leaks.
47. Ref. 34, Section 15-0.
48. Personal Communication. Nagy, D., BS&B Safety Systems,
Inc., Charlotte, NC, with D. J. Whitt, Radian
Corporation. October 30, 1990.
49. Personal communication. Moretti, E., Radian Corporation,
with Ponder, T., PEI Associates, Dallas, TX. February 2,
1990 and November 22, 1990.
3-64
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50. Personal communication. Moretti, E., Radian Corporation
with Darbonne Services, Inc., Lake Charles, LA.
November 27, 1989.
51. Telecon. Whitt, D.J., Radian Corporation, with
Johnson, J., Foxboro Equipment Company. October 26,
1990. Pricing of Foxboro equipment.
52. Langley, G.J., S.M. Dennis, L.P. Provost, J.F. Ward.
(Radian Corporation). Analysis of SOCMI VOC Fugitive
Emissions Data. Prepared for U.S. Environmental
Protection Agency. Research Triangle Park, North
Carolina. Publication No. EPA-600/2-81-111. June 1981.
53. Memorandum from Moretti, E., Radian Corporation, to
Markwordt, D., EPA/CPB. May 22, 1989. Summary of Work
Assignment 42 - Equipment Leak Emission Estimates,
VOC/HAP.
3-65
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APPENDIX A
EXAMPLE COSTS FOR INSTALLATION OF A FLARE TO A PROCESS VENT
This is an example calculation to determine the total
annual costs associated with controlling an example process
vent stream with a flare having a removal efficiency of
98 percent.
The purpose of this appendix is to demonstrate the
approach used in the HON analysis. In the calculations below,
all significant figures have been retained until the final
calculation to make it easier for the reader to follow the
calculation and to avoid potential error due to round off of
intermediate calculations. It should not be inferred that the
intermediate results represent the actual number of
significant figures.
The example stream has a volumetric flow rate (Qtot) °f
7.39 scf/min and contains ethylbenzene. Other design and
costing assumptions are presented in Tables A-l and A-2. The
design and costing procedure follows the one set forth in the
Office of Air Quality Planning and Standards Control Cost
Manual (OCCM)1 and described in Section 3.1.1 of the text of
this volume.
1. Determine the auxiliary fuel requirement necessary to
meet the Federal requirement that a stream have a minimum net
heating value (Bv) of 300 Btu/scf, as stated in 40 CFR
Section 60.18 of the Federal Register. The example stream has
a heating value of 321 Btu/scf, so no auxiliary fuel is
required.
2. Calculate the minimum flare tip diameter (Dmin) based
on the maximum stream exit velocity. For a stream with a heat
A-l
-------
TABLE A-l. DESIGN ASSUMPTIONS
Parameter Description Value
Flare type stream assisted, elevated
Minimum heat content of 300 Btu/scf
combusted stream (vent + fuel)
Minimum flare tip diameter 1 in.
Minimum flare height 30 ft.
Maximum flare radiation level 500 Btu/hr-ft2
Fraction of combustion heat 0.2
radiated
Fraction of radiated heat 1.0
transmitted
Minimum purge gas velocity 0.04 ft/sec
Steam to waste gas ratio 0.4 Ib steam/Ib waste gas
Minimum knockout drum diameter 6 in.
A-2
-------
TABLE A-2. COST ASSUMPTIONS
Cost Item
Value
Direct Annual Costs
Dollar basis
Operating labor
Operator
Supervisor
Labor cost
Maintenance
Labor
Materials
Labor cost
Utilities
Electricity
Purge gas
Pilot gas
Auxiliary fuel
Steam
Electricity cost
Fuel cost
Steam cost
Indirect Annual Cost
Overhead
Administrative charges
Property tax
Insurance
Capital recovery5
July, 1989 (CE Index = 355.9)
630 hr/yr
15% of operator
$13.20/hr
1/2 hr/shift
100% of maintenance labor
$14.50/hr
All utilities equal to
(consumption rate) x (hrs/yr) x
(unit cost)
$0.0509/kW»hr
$3.03/1000 scf
$7.77/1000 scf
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
aThe capital recovery factor is based on the assumption of an
equipment life of 15 years and an interest rate of 10% with
no inflation over the equipment life.
A-3
-------
value greater than 300 Btu/scf and less than 1000 Btu/scf, the
maximum gas velocity, Vmax (ft/sec) , is given by:
1°910 (Vmax) = (Bv + 1214J/852
Vmax = 10 t(Bv + 1214)/852]
10[ (321 + 1214)/852]
10 C1-8) =63.3 ft/sec
The minimum flare tip diameter (Dmin) is then given by:
Dmin(in, . 1.95 "tot <»3/min,
N vmax (ft/sec)
Where:
Qtot = vent flow rate + auxiliary fuel flow rate,
=1-95
7.39 ft3/min
63.33 ft/sec
= 0.67 in
The minimum flare tip diameter is 1 in, so the rest of the
design and costing will be based on a 1-in diameter flare tip.
3. Next, the flare height L, is calculated using:
47TK
where :
rj = Fraction of heat intensity transmitted;
f = Fraction of heat radiated;
R = Net heat release (Btu/hr) ;
K = Allowable radiation (500 Btu/hr»ft2); and
R( Btu/hr) = Bv X Qtot x 60 min/hr.
In order to produce a conservative design, assume that all the
heat produced is transmitted (r? = 1) and the fraction of heat
radiated, f, is 0.2.
A-4
-------
R = (7.39 scfm)
60
mm
hr
321
Btu
scf
= 142331
Btu
hr
and
L =
(1.0) (0.2)
142331
Btu
hr
4 n 500
Btu
hr-ft2
0.5
= 2.13 ft,
The minimum design flare height is 30 ft. so costing will be
based on a 30 ft. flare.
4. Purge gas is used to keep a minimum required positive
flow throughout the system. The purge gas requirement (FpU)
is estimated according to flare tip diameter by:
Fpu (scf/yr) = 6.88 x io3 D2(in)
•pu
scf
yr
V J /
TO -
6.88 * 10J (I)2 = 6.88 * 10
scf purge gas
5. Pilot gas is needed to keep the burners lit. The
pilot gas requirement (Fpi) is calculated by:
Fpi (scf/yr) = (6.13 x 105) N
where:
N = Number of pilot burners (see table below).
Flare Tip Diameter fin)
1-10
12-24
30-60
>60
Number of Pilot Burners (N)
1
2
3
4
The flare tip diameter is 1 in., so the number of pilot
burners, N, is 1.
A-5
-------
pi
scf
yr J
= 6.13 *
(1) - 6.13
yr
6. A steam-assisted flare uses steam to create
turbulence for even mixing of the vent stream in the burner
flame. The amount of steam required is based on the mass flow
of the vent stream (W). The ideal gas law is used to convert
the volumetric vent stream flow to mass flow.
W (Ib/hr) = Qvent (scfm) * 60 (min/hr) * P/RT * MWavg
where:
= Stream pressure (atm);
= Steam temperature (°R);
gas constant = 0.7302 ft3 * atm/lb-mol*°R;
= average molecular weight of stream (assume
29.0 Ib/lb-mol).
P
T
R
MW
avg
W
Ib
hr
(7.39 scfm)
/
0.7302
Imin
fin
hr J
(1 atm)
atm- ft3
lb-mol-°R
Ib
TO
lb-mole,
(110.7 + 460)°R
W = 30.8 Ib/hr
The steam requirement formula in the OCCM is
S (Ib/yr) = 3500 * W
For the example stream,
S (Ib/yr) = 3500 (30.8) = 1.08 X 105 Ib/yr
7. A knock-out drum is needed to eliminate mist
entrainment in the vent stream. For the HON analysis, the
densities of the pollutant vapor and liquid phases are
estimated using propane (€3^). Propane was selected because
it has a similar molecular structure to that of the design
molecule (C2.85H5.7°0.63) described in Section 3.1.2.1.1 of
the text.1'2 The vapor stream density, pv, is 0.111 Ib/ft3,
A-6
-------
and the condensed liquid density, p£, is 37 Ib/ft3. The
maximum linear velocity of the stream, U, is calculated by:
U (ft/sec) = G [(p£-pv)/pv]0-5
The factor G is assumed to be 0.2 as suggested on page 7-26 of
the OCCM. For the example stream,
U (ft/sec) = 0.2 [(37-0.111)/0.Ill]0-5 = 3.65 ft/sec
The cross sectional area of the drum, A, is calculated as the
volumetric flow rate of the vent stream divided by the maximum
linear velocity of the stream.
A (ft2) = Qvent (scfm)/[60 (sec/min) * U (ft/sec)]
A = 7.39 ft3/min/(60 sec/min * 3.65 ft/sec)
A = 0.034 ft2
The minimum drum diameter, tmin, is:
(in) = 13.5 * A (ft2) 0.5
= 13-5 (0.034)0-5 = 2.5 in
The minimum available drum size is 6 inches according to the
OCCM. For drums of diameter less than 36 in., the drum wall
thickness is 1/4 in. The drum height is typically three times
the diameter. In this case, the drum height would be
3(6 in) = 18 inches.
8. The three main components of capital cost for flares
are the flare itself, the knockout drum, and the piping for
the vent stream. The flare cost, Cp, equation varies with
flare height and diameter. For flares of height less than 100
feet the equation is:
CF (March, 1990 $) = (78.0 + 9.14D + 0.749L)2
For this example the flare cost is:
CF = [78.0 + 9.14(1.0) + 0.749(30)]2 (355.9/354.6)
$12,058
when adjusted to July, 1989 dollars.
9. The knockout drum cost, C^, is dependent upon the
height (h), diameter (d), and wall thickness (t) of the drum:
CK (March, 1990 $) = 14.2 [d * t * (h + 0.812d)]0.737
The cost of the example drum is:
CK = 14.2 [(6)(0.25)(18 + 0.812(6))]0.737 (355.9/354.6)
= $193 (July, 1989 dollars)
A-7
-------
10. The piping cost, Cp, depends on the diameter and
length of pipe required. The pipe diameter is assumed to
equal to the flare tip diameter, and the piping length
required is assumed to be 300 feet. The piping cost equation
is:
Cp (March, 1990 $/100 ft) = 1.27D1-21
For the example stream,
Cp = (300 ft.) (1.27) (l.O)1-21 (355.9/354.6)
= $381 (July, 1989 dollars)
11. The total equipment cost (EC) is the sum of the
flare cost, the knockout drum cost, and the piping cost.
EC = Cf + Ck + Cp = $12,058 + $193 + $381
= $12,632 (July, 1989 dollars)
12. Factors found in the OCCM allow the calculation of
purchased equipment cost (PEC) then total capital investment
(TCI).
The equations are:
PEC = 1.18 EC = 1.18 ($12,632) = $14,906
and
TCI = 1.92 PEC = 1.92 ($14,906) = $28,619
13. Direct annual costs include labor, materials, and
utilities. Labor and materials costs are calculated as a
function of hours of labor required and are shown in
Table A-3. Utilities costs are calculated from the product of
consumption rate and unit cost. The annual consumption values
appearing in Table A-3 for the utilities are calculated as
follows:
Annual Gas Consumption = Auxiliary Gas + Purge Gas + Pilot Gas
0 + (6.88 * 103 scf/yr) + (6.13 * 105 scf/yr)
619880 scf/yr
Annual Steam Consumption = 1.08 * 105 Ib/yr = 108,000 Ib/yr
A-8
-------
TABLE A-3. ESTIMATION OF TOTAL ANNUAL COST FOR A FLARE SYSTEM
Cost Component
Cost Factor
Annual
Consumption
Annual
Cost3
Direct Annual Costs
Utilities
Gas
Steam
Electricity*3
Labor
Operating Labor0
Supervision &
Administration
Maintenance
Labord
Materials
$3.03/1000 scf
$7.77/2000 Ib
$0.0509/kW-hr
619880 scf
108,000 Ib
1 kW-hr
$13.20/hr 630 hrs
15% of Operating
Labor
$14.50/hr
100% of
Maintenance
Labor
547.5 hrs
$1,878
420
0
$8,316
$1,247
$7,939
$7,939
TOTAL DIRECT ANNUAL COST (TDAC)
$27,739
Indirect Annual Costs
Overhead
General &
Administrative
Property Taxes
Insurance
Administrative
Charges
60% of all labor
and materials
1% of TCI
1% of TCI
2% of TCI
Capital Recovery6 13.15% of TCI
$15,265
$286
$286
$572
$3,763
TOTAL INDIRECT ANNUAL COST (TIAC)
$20,172
TOTAL ANNUAL COST
(TAG)
TDAC + TIAC
$47,911
aJuly 1989 dollars
^Assumed default value of 1 kW-hr (see Step 13).
GAssumed flare system operated continuously and required 630
hours of labor per year.
^Assumed flare system operated continuously and required
0.5 hour of labor per 8-hour shift.
eSee Table A-2.
A-9
-------
Annual Electricity Consumption = (1.17 * 10-4)
* (Qvent scfm) (Ap) 10.6
Ap (in H20) = [(1.238 * 10~6) (QVent) ~ i-15 * 10~4]
* (pipe length)
Ap = [(1.238 * 10~6) (7.39) - 1.15 * 10~4] (300 ft)
= -0.032 inches H20
The above pressure drop correlation was developed for use in
the RON costing procedure; however, the lower limit of its
applicability is a vent stream flow of 93 scfm since a
negative pressure drop is not technically realistic. For vent
stream flow rates below 93 scfm, the subsequent calculation of
power consumption is set to a default value of 1 kW«hr/yr.
This default value was used for the example process vent
because its vent stream flow rate is only 7.39 scfm.
14. Indirect annual costs include overhead, taxes,
insurance, and administration costs. Factors and costs for
each of these appear in Table A-3.
15. Total annual cost is the sum of the direct annual
cost and the indirect annual cost. The total annual cost for
the application of a flare to this example stream is
approximately $47,900.
A-10
-------
REFERENCES
1. Blackburn, J.W. Control Device Evaluation Thermal
Oxidation. In: Organic Chemical Manufacturing, Volume
4: Combustion Control Devices. Prepared for U.S.
Environmental Protection Agency, Research Triangle Park,
NC. December 1980. Report 1. pages 111-2,5.
2. Memorandum from Scott, K., Radian Corporation, to RON
project file. February 5, 1992. Composition of design
molecule.
A-ll
-------
APPENDIX B
EXAMPLE COSTS FOR INSTALLATION
OF AN INCINERATOR/SCRUBBER
SYSTEM ON A TRANSFER LOADING RACK
The following example illustrates the procedure used to
calculate the total annual costs and the total capital costs
associated with the addition of an incinerator plus scrubber
system to a tank truck or tank car transfer rack. The purpose
of this appendix is to demonstrate the approach used in the
HON analysis. In the calculations below, all significant
figures have been retained until the final calculation to make
it easier for the reader to -.follow the calculation and to
avoid potential error due to round off of intermediate
calculations. It should not be inferred that the intermediate
results represent the actual number of significant figures.
The example tank truck transfer rack is located at an
example facility which loads ethylene dichloride,
formaldehyde, inethanol, and vinyl chloride. The costing
procedure for an incinerator and scrubber assigned to a tank
car transfer rack is identical to the procedure presented
below. Calculation data are provided in Table B-l for the
facility.
B.I TOTAL CAPITAL COSTS
The total capital cost of the transfer rack control
device includes the capital cost of the incinerator and the
scrubber. For the purpose of the HON analysis, it was assumed
that a nitrogen blanketing system, a vapor collection system,
and a submerged loading system were already in place at each
facility for each vehicle. Therefore, there were no capital
costs associated with these three systems.
B-l
-------
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-------
B.I.I Incinerator Capital Investment
The incinerator capital cost includes the cost of the
incinerator equipment and any associated ducts and fans. The
cost of the incinerator is directly dependent on the total
volumetric flow rate through the control device.
1. Calculation of the amount of VOC entering the
incinerator. The amount of VOC entering the
incinerator is calculated as follows:
Mols VOC In
(Ib-mol/hr):
. . Mg
Baseline Emissions —
Mg
1 Ib
453.593 g
1 yr
Operating Hours
Molecular Weight of VOC
Ib
Ib-mol
The baseline emissions of VOC are the emissions generated
during the vehicle loading operation which are routed to the
incinerator. The operating hours represent the number of
hours per year that the control device will operate, and the
molecular weight is representative of the VOC molecules
entering the incinerator.
la. Calculation of baseline emissions.
To calculate the baseline emissions associated with tank
car and tank truck transfer racks, the uncontrolled emissions
corresponding to each material must first be calculated. This
is done for each vehicle by multiplying the actual throughput
of the material by the emission factor of the material. If
the loading of a material is regulated by State or Federal
control, the uncontrolled emissions for each material are
reduced by the quantity (1 - the fractional control level
efficiency) to obtain baseline emissions. If the loading of a
material is not controlled at the State or Federal level, the
baseline emissions are equal to the uncontrolled emissions.
B-3
-------
The following equations illustrate the calculation of the tank
truck baseline emissions for the methanol produced at the
example facility.
The tank truck emission factor for methanol is calculated
from the following equation:1
Emission Factor (lb/1000 gal) = 12.46 * S * VP * MW/(T + 460)
where:
S = 0.6 = A saturation factor for submerged loading;
VP = 1.93 = Vapor pressure of methanol (psi);
MW = 32.04 = Molecular weight of methanol (Ib/lb-mol);
T = 77 = Loading temperature (°F)
Ib
Emission . Ib-mol Ib
= 12.46*0.6*1.93 psi*32.04 =0.861
Factor 537 1000 gal
Therefore:
/ Actual gal
Throughput
/Emission
I
Ib
1000 gal
Tank Truck
Uncontrolled =
Emissions
6.85*10
gal
yr J
0.8611b
103 gal
453.593
Ib
Mg
106 g
= 2.68
Mg
yr
Assume the example facility is located in a State which
requires 90 percent control of emissions from materials having
a vapor pressure greater than or equal to 1.5 psi and an
actual throughput to the tank truck rack greater than
40,000 gal/day. The throughput of a chemical in gallons per
day is calculated by dividing the annual throughput by the
number of days per year that a chemical is transferred. It is
assumed that facilities transfer a chemical for the same
number of days per year that they produce the chemical.
Multiplying the number of days in a year by the capacity
B-4
-------
utilization factor results in an estimate of the number of
days in a year that the facility transfers product. The
capacity utilization factor represents the difference between
the amount of product actually produced and the maximum
capacity. Of the four chemicals produced and transferred at
the facility, ethylene dichloride and methanol meet the vapor
pressure criteria for State control. However, none of the
chemicals meet the tank truck throughput requirement for State
control. Therefore, the baseline emissions for methanol and
ethylene dichloride transfer are equal to the respective
uncontrolled emissions because they do not meet the
requirements for State control. The emission factors for
vinyl chloride and formaldehyde are zero because these
materials have a vapor pressure greater than 14.7 psi.
Although there is a NESHAP for vinyl chloride, it is assumed
that vinyl chloride and formaldehyde are transferred under
pressure and their uncontrolled and baseline emissions are
equal to zero. Therefore, there are no control costs
associated with a chemical having a vapor pressure greater
than atmospheric pressure. The data and results are shown in
Table B-l.
To obtain the total baseline emissions for the tank truck
transfer rack, the baseline emissions from tank truck loading
operations for each material are summed. For the example
facility, the tank truck rack total baseline emissions are
4.55 Mg/yr.
Ib. Calculation of operating hours.
Operating hours for the control device are calculated by
doubling the loading time. This allows for warm-up of the
incinerator prior to loading and cool-down after loading is
completed. Loading time is calculated by dividing the total
annual rack actual throughput by the total filling rate. The
total rack actual throughput is calculated by summing the
actual rack throughput of every transferred chemical. The
total filling rate is equal to the vehicle filling rate
multiplied by the number of arms per rack. In the HON
analysis, it is assumed that the vehicle filling rates of tank
B-5
-------
trucks and tank cars are 170 gpm and 190 gpm, respectively.
The number of arms per transfer rack are determined by the
maximum total annual rack throughput and the total number of
materials transferred as shown in Tables B-2 and B-3.
For the example facility, the total annual maximum
throughput to the tank truck transfer rack is 15.26 MMgal/yr.
Since this facility transfers four materials, it is assigned
one 4-arm transfer rack. Therefore, the operating hours for
the example facility are calculated from the total annual
actual throughput as follows.
Operating
Hours =
(hr/yr)
12.53
gal
170
min- arm
f- gal
y T 0°
yr
60 min
y * '
hr
I arms
* 2 = 614.2
hr
yr
Ic. Calculation of molecular weight.
The molecular weight of the stream sent to the
incinerator is a weighted average of the molecular weights of
the individual chemicals in the stream. To obtain a rack
molecular weight, the molecular weight of each transferred
chemical is weighted based on the quantity of the chemical
sent to the incinerator. The quantity of the chemical routed
to the incinerator is equal to the baseline emissions of the
chemical. The following equations are used to calculate the
number of Ib-moles of ethylene dichloride and methanol
entering the incinerator. Formaldehyde and vinyl chloride are
not included in the molecular weight calculation, because it
is assumed that the two chemicals have no baseline emissions
(see Step la).
Ethylene
Dichloride
(1.874 X 106 g/yr)
Ib-mol 1
, Yr J
454 —
( IbJ
/
Ib
t QQ Qf.
\ Ib-mol
= 41.71
Ib-mol
yr
B-6
-------
TABLE B-2. MODEL TANK CAR TRANSFER RACKS
Number of
Materials
1-3
4-9
10 - 22
> 23
Throughput (TP) Range
(MMgal/yr)
0 < TP < 10
10 < TP < 40
40 < TP < 80a
0 < TP <10
10 < TP < 20
20 < TP < 30
30 < TP < 60b
0 < TP < 3
3 < TP < 80C
0 < TP < 10
10 < TP < 20d
Number of
Arms
3
8
16
3
6
10
16
3
10
4
9
aFor throughputs above the maximum value, add an additional
3-arm rack per 10 * 106 gal.
bFor throughputs above the maximum value, add an additional
3-arm rack per 10 * 106 gal.
°For throughputs above the maximum value, add an additional
3-arm rack per 3 * 106 gal.
dFor throughputs above the maximum value, add an additional
4-arm rack per 10 * 106 gal.
B-7
-------
TABLE B-3. MODEL TANK TRUCK TRANSFER RACKS
Number of Throughput (TP) Range
Materials (106 gal/yr)
1-4 0 21 0 < TP < 12
12 < TP < 24d
Number
of
Arms
1
2
4
1
2
4
6
1
3
4
aFor throughputs above the maximum value, add an additional
1-arm rack per 3 * 106 gal.
bFor throughputs above the maximum value, add an additional
1-arm rack per 3.5 * 106 gal.
°For throughputs above the maximum value, add an additional
1-arm rack per 15 * 106 gal.
dFor throughputs above the maximum value, add an additional
4-arm rack per 12 * 106 gal.
B-8
-------
Methanol
/ \
J-D HIOJ.
yr
\ •* /
(2.680 X 106 g/yr)
/ \
454 —
IbJ
*
\
Ib
3° 04
Ib-mol J
= 184.24
Ib-mol
yr
Tank Truck
Rack
Molecular
Weight =
Ib
Ib-mol
41.71
Ib-mol
yr
41.71 + 184.24-
Ib-mol
yr
* 98.96
Ib
184.24
Ib-mol
yr
41.71 + 184.24
= 44.39
Ib
Ib-mol
Ib-mol
Ib-mol
yr )
* 32.04
Ib
Ib-mol
Therefore, to calculate the amount of VOC entering the
incinerator for the tank truck transfer rack at the example
facility:
Mols VOC In =
Mg
55—-
• *J *J ^^^^^
yr)
^ f 106 g 1 Ib
1 Mg 1 I 453.593 g
V -* s V -* /
•.
1 yr
614.2 hrl
+
J
/ \
1
Ib
\ \ 39
Ib-mol
= 0.368
Ib-mol
hr
B-9
-------
2. Calculate the air required to combust the VOC
molecules.
The combustion air requirements are based on a model VOC
molecule, C2.85H5.7°0.63 as explained in Section 3.1.2.1.1 of
the text. Therefore, the combustion mole balance reads as
follows:
C2.85H5.7°0.63 + 3-96 °2 —> 2-85 C02 + 2.85 H20
There are 3.96 moles of oxygen required to combust 1 mole of
VOC. Assuming air consists of 20.9 mole percent oxygen, the
combustion air requirements for the example facility are:
Combustion Ib-mol
Air
3.96
Ib-mol O-> 1 Ib-mol Air
*-• I ~tf I
Ib-mol VOCj (0.209 Ib-mol O2
= |o.368 VOC
hr
Ib-mol Air
= 6.97
hr
3. Calculation of dilution air.
The incinerator exhaust is required to contain at least
3 mole percent oxygen. Since all of the oxygen in the
combustion air is reacted in the incinerator, a certain amount
of dilution air must be fed to the incinerator. Using a total
balance around the incinerator, it is found that the
volumetric flow of exhaust is equal to the sum of the
volumetric flow of dilution air, combustion air, and VOC
entering the incinerator. If it is assumed that all of the
combustion oxygen is reacted in the incinerator, and that
volume percent is approximately equal to mole percent at
atmospheric pressure, then the volumetric flow rate of the
dilution air is calculated as follows:
(Flow Rate ofl (Flow Rate of
0.209 * L--, .. „ . =0.03 _ ,
Dilution Air Exhaust
Exhaust (scfm) = VOC (scfm) + Combustion Air (scfm) +
Dilution Air (scfm)
B-10
-------
For the tank truck rack at the example facility,
Combustion
Air
(scfm)
Ib-mol
6.97 *
hr
=45.5 scfm
1 hr
60 min
392 scf
Ib-mol
VOC Stream
(scfm)
Total filling rate
170-
gal
mm. arm
90.91 scfm
* (4 arms) *
1 ft3
7.48 gal
(Dilution) (90.91 scfm + 45.5 scfm)
0.209 * , . =0.03 * -,--, 4.-
^ Air J I + Dilution Air J
Dilution Air = 22.87 scfm
4. Calculate the auxiliary fuel requirement.
Auxiliary fuel is required for startup of the incinerator
unit and is also required to maintain the reactor temperature
and to stabilize the flame. Auxiliary fuel is also needed to
sustain combustion if the heat content of the VOC stream is
low. For the HON analysis, the energy provided by the
auxiliary fuel must be at least five percent of the total
energy input to the incinerator.
4a. Calculate the volumetric flow rate of the auxiliary
fuel.
The volumetric flow rate of auxiliary fuel is calculated
using the following equation which appears in the OAQPS
Control Cost Manual.2
Qaf pf i Qf i[cpair (1.1 Tfi-Tfo- .1 Tref) - (-A hcfo)]
(scfm) =
paf
(Tfi_Tref)]
B-ll
-------
where:
pfi = density of the inlet gas (lb/ft3)
Qfi = flow rate of the inlet gas (ft3/min)
cpair = specific heat of air (Btu/lb °F)
Tfi = operating temperature of the incinerator
( F)
Tfo = temperature of the outlet gas (°F)
Tref = reference temperature (°F)
Ahcf0 = heat of combustion of outlet gas (Btu/lb)
Ahcaf = heat of combustion of auxiliary fuel
(Btu/lb)
paf = density of the auxiliary fuel (lb/ft3)
Because the inlet and outlet gas streams consist mainly of
air, it is assumed that the density of these streams is equal
to the density of air. The operating temperature of the
incinerator is 1600 °F if the inlet stream is nonhalogenated
and 2000 °F if the inlet stream is halogenated. The reference
temperature is 77 °F, and it is assumed that the temperature
of the outlet gas stream is also 77 °F. It is also assumed
that the heat content of the outlet stream is approximately
equal to the heat content of the inlet stream, because these
streams consist mainly of air. The auxiliary fuel is methane.
The heat content of VOC stream is calculated and then
used to find the total energy input. The heat content of the
VOC generated by the loading operation is the sum of the
respective heat contents of each transferred chemical divided
by the operating hours and the VOC stream flow rate.
B-12
-------
Heat Content
(Btu/scf)
Chemical
Baseline
Emissions g/yr,
Chemical
Heat
of Combustion
Btu/lb
VOC Stream
Flow Rate
SCFM
Operating
Hours
hr/yr _
60
mm
hr
* 453.593
Ib
For the example facility, the VOC stream contains ethylene
dichloride and methanol.
Heat Content _ [(1.874 *106) * (3400.0) + (2. 679 *106) * (8419)]
(Btu/scf) (614.2) * (90.91) * (60) * (453.593)
= 19.03
Btu
scf
The heat content of the VOC stream is then adjusted to
incorporate the lower heat content of the combustion and
dilution air fed to the incinerator with the VOC. The total
inlet gas flow rate is equal to the sum of the flow rates of
the combustion air, the dilution air, and the VOC stream.
New
Heat Content =
(Btu/scf)
Heat Content
of VOC Stream
Btu
scf
VOC Stream Flow Rate (scfm)
Total Inlet Gas
Flow Rate (scfm)
Btu 90.91 scfm Btu
= 19.03 * = 10.86
scf (159.28 scfmj scf
The new heat content represents the heat content of the
incinerator inlet stream and is therefore approximately equal
to the heat content of the outlet stream. If the new heat
content is still larger than 98 Btu/scf for nonhalogenated
streams or larger than 95 Btu/scf for halogenated streams,
more dilution air is required to prevent exceeding design
specifications. The heat content is then converted to units
of Btu/lb.
B-13
-------
Ahcfo =10.86
Btu
scf
1 scf
0.0739 Ib
= 146.96
Btu
Ib
For the example facility, the following input data were
used to calculate the flow rate of the auxiliary fuel. Since
the example facility transfers a halogenated chemical,
ethylene dichloride, the incinerator operates at 2000 °F.
pfi = pair = 0.0739 lb/ft3
Qfi = VOC Stream + Combustion Air + Dilution Air
90.91 scfm + 45.5 scfm + 22.87 scfm
159.28 scfm
cpair = 0.255 Btu/lb °F
Tfi = 2000 °F
Tfo = 77 °F
Tref = 77 °F
AhCf0 = -146.96 Btu/lb
Ahcaf = -21,502 Btu/lb
paf = 0.0408 lb/ft3
(0.0739) * (159.28) *[0.255*(!.!* (2000) -77 - (0.1* 77) -146.96]
Qaf=
(0.0408) *[21,502-l.l*(0.255*(2000-77))]
= 5.40ft3/min
4b. Calculate the air required to combust the auxiliary
fuel.
CH4 + 202 > CO2 + 2H20
To combust one mole of methane, two moles of oxygen are
required. It is assumed that air consists of 20.9 mole
percent oxygen.
B-14
-------
Air Required
to Combust
Methane
(scfm)
2 mol 02
mol CH4
1 mol Air
0.209 mol O2,
* (5.40 scfm CH4) *
0.0408
lb
CH/
scf
= 51.65 scfm
1 Ib-mol
16 lb CH4,
392 scf
1 Ib-mol Air
4c. Calculate the total energy input.
The total energy input is the energy contained in the VOC
stream entering the incinerator.
TEI = pfi * Qfi * (Tfi - Tref) * Cpfi
The specific heat of the outlet stream, Cpfi, is approximately
egual to the specific heat of air. The inlet stream flow
rate, Qfi, will now include the flow rate of the auxiliary
fuel and the flow rate of the air required to combust the
auxiliary fuel. The total energy input for the example
facility is shown below.
TEI =
0.0739
lb
scf
* (l59.28 scfm + 5.40 scfm + 51.65 scfm)
* 2000°F - 77°F *
0.255-
Btu
lb°F
= 7839.4
Btu
4d. Calculate the auxiliary fuel energy input.
The auxiliary fuel energy input equals the energy
contained in the methane entering the incinerator.
AFEI - paf * Qaf * (-Ahcaf)
where:
paf =
Qaf =
Ahcaf
density of methane (lb/ft3)
volumetric flow rate of methane (scfm)
= heat of combustion of methane (Btu/lb)
B-15
-------
For the example facility,
AFEI -
0.0408
Ib
* (s.40 scfm) *
21,502
Btu
= 4737.3
Ib
Btu
min
The energy input of the auxiliary fuel must be greater
than 5 percent of the total energy input.
Btu Btu
TEI * 0.05 = 7839.4 * 0.05 = 392
min min
The auxiliary fuel energy input is much greater than
392 Btu/min.
5. Calculate the capital cost of the incinerator.
The cost of the incinerator is based on the total
volumetric flow to the control device. The total inlet
volumetric flow is equal to the sum of the flow rates of the
VOC stream, the combustion and dilution air, and the auxiliary
fuel and associated combustion air. For the example facility,
the total volumetric flow to the incinerator is equal to 216.3
scfm. However, in the HON analysis, it is assumed that the
smallest incinerator is designed for a flow of 500 scfm.
Therefore, the equipment cost of any incinerator with a flow
less than 500 scfm will be based on a flow of 500 scfm.
The equipment cost equation is based on zero percent heat
recovery, because the stream is halogenated. Cost equations
for the incinerator are given in Table 3.1-5 of the text.
/
n o-mc: Price Index
EC = 10294 * (Qtot)°-23b5 *
(^ $340.1
The price index will adjust the cost of the incinerator to
July, 1989 dollars.
B-16
-------
n 9-^R $355.9
EC = 10294 * (500)0'2355 *
$340.1
= 46,549.30
It is assumed that the ducting associated with the incinerator
is 300 feet long with two elbows per 100 feet. The ducting
also has a diameter of 24 inches and is made of 1/8 inch thick
carbon steel. The ducting cost is also adjusted to July, 1989
dollars.
TTf
= [210 * (24) °'839+2 *4.52 * (24) 1'43]*3
$355.9
* = $11,732.41
$352.4
The cost of the fan is proportional to the volumetric flow
rate of the VOC which the fan must send to the incinerator.
The cost is adjusted to July, 1989 dollars.
EC(fan) = (96.96418 * (VOC Flowrate SCFM)°-547169)
Cost Index
*
$342.5
Ec(fan) = (96.96418 * (90.91 scfm)°•547169)
$355.9
$342.5
= $1188.42
The basic equipment cost (EEC) is equal to the cost of the
incinerator and the associated fans and ducting.
EEC = $46,549.30 + $11,732.41 + $1188.42 = $59,470.13
The purchased equipment cost (PEC) includes the EEC and the
cost of instrumentation, sales tax, and freight.
B-17
-------
Instrumentation =0.1* EEC
Sales Tax = 0.03 * EEC
Freight = 0.05 * EEC
PEC = 1.18 * EEC
PEC = 1.18 * 59,470.13 = $70,174.75
The total capital investment is calculated by multiplying the
purchased equipment cost by an installation factor of 1.25.
TCI = 1.25 * PEC
TCI = 1.25 * $70,174.75 = $87,718.44
B.2.1 Scrubber Capital Investment
For the HON analysis, the scrubber design and capital
investment is based on a packed bed, countercurrent scrubber
tower. The scrubber capital investment includes the cost of
the scrubber, packing, platform, stack, and any associated
ducts and fans. The capital investment is directly related to
the size of the scrubber and the pressure drop across the
scrubber, which is dependent on the liquid and vapor flow
rates through the column. The absorbing liquid is water, and
the vapor consists mainly of air. The design and costing
procedures closely follow the procedure presented in Chapter 9
of the OCCM (Gas Absorbers) , and the Handbook on Control
4
Technologies for Hazardous Air Pollutants (HAP Manual) .
1. Calculate the flow rate of acid gas to the scrubber.
The amount of vapor flowing to the scrubber is equal to
the amount of vapor leaving the incinerator. The vapor stream
is predominantly air.
Mols Vapor
Ib-mol
= (216.3 scfm) *
hr
Ib-mol
60
mm
hr
1 Ib-mol
392 scf (at T = 77°F)
= 33.11
hr
2. Calculate the liquid flow rate through the scrubber.
It is assumed that the liquid and vapor flow rates
through the column are essentially constant. Therefore, the
B-18
-------
liquid to vapor flow rate ratio throughout the column is also
constant and, for the purpose of the HON analysis, equal to
17 gpm/1000 scfm. By converting both quantities to Ib-moles
per hour and multiplying the ratio by the vapor flow rate, the
liquid flow rate is found.
17 gpm
1000 scfm
gal
17
min
v /
100
* 60
n
min
mir
hr
*
-)'
y
8.34
V
1 lb-mol
392 scf
lbl.
galj
*
\
lb-mol
18 Ib
/
min
60
hr
472.6
lb-mol
hr
water
lb-mol
153.06 vapor
hr
Liquid
Flow Rate
Through
Scrubber
(Ib-mol/hr)
472.6
lb-mol
water
hr
153.06
lb-mol
vapor
hr
* 33.11
lb-mol Vapor
hr
= 102.23
lb-mol
hr
3. Calculate the diameter of the tower.
As explained in Section 3.1.2.1.5 of the text, a
correlation for randomly packed towers based on flooding
considerations is used to determine the tower diameter. The
vapor flow rate per cross-sectional area of the column is
first determined by the correlation. By knowing the vapor
flow rate, the cross-sectional area of the column can then be
found.
B-19
-------
ABSCISSA =
i lb
Liquid Flowrate |
hr
Ilb
hr
Density of Vapor (pv)
Density of Liquid (pL/
.5
Ib-mo
in0 T\
\ hr
/ \
lb~mol
n 1 1
hr
1
j
^
*
tf
1
18 lb
Ib-mol
V s
lb
lb-mo:
\
L
*
lb
0 0719
ft3
lb
fi9 o
I f t J J
= 0.0661
.5
ORDINATE = -0.9809237 * (ABSCISSA)(~°•0065226 * In (ABSCISSA))
+ (ABSCISSA)(-0.021897)
= 0.126
The vapor flow rate per cross-sectional area is calculated by
the following equation. For the purpose of the HON analysis,
the column is assumed to operate at 60 percent of flooding.
Vapor Flowrate
Area
s-Sectio
lb
sec- ft2
nal
ORDINATE * pv * pL * gc
H
[ U3J
f HL
(2.42J
.2
1/2
* f
where:
pv
PL
9c
HL/2.42
density of the vapor = 0.0739 lb/ft3
density of water =62.2 lb/ft3
gravitational constant =
32.2 ft»lbm/lbf»sec2
viscosity of solvent = 0.85 cp
B-20
-------
a/e3
the void space of packing per surface area
to volume ratio of packing = 69.1 ft
flooding factor = 0.6
Vapor
Flowrate
per Cross-
Sectional
Area
Ib
sec
•ft2
32.2
0.126*0.0739 * 62. 2 * *
69.1 (0.85)0'2
Vapor Flowrate
per Cross-Sectional = 0.3168
Area
Ib
sec- ft'
Area of Tower
(ft2)
Vapor Flowrate j
sec;
Vapor Flowrate per
Cross-Sectional Area
Area of Tower
(ft2)
33.11
Ib mol\
hr
* 29
Ib \
Ihr
Ibmol/ \3600 sec
0.3168
Ib
sec-f t2
= 0.842 ft2
A - TC * (radius) 2
* «Diameter \2
Diameter
(ft)
Diameter - /— * 0.842\2 = 1.035 ft
4. Determine the number of transfer units.
The number of transfer units represent the number of
theoretical equilibrium stages required to absorb the vapor
pollutant. This number is determined from a quantity called
the absorption factor which is calculated by dividing the
B-21
-------
liquid to vapor ratio by the slope of the equilibrium line.
The slope of the equilibrium line is represented by the
difference in the vapor mole fractions divided by the
difference in the liquid mole fractions from the top to the
bottom of the column. For the purpose of the RON analysis,
the slope is equal to 0.1.
Absorption
Factor
472.6 Ib-mol/hr water *
18 Ib
Ib-mol
153.06 Ib-mol/hr vapor *
29 Ib
Ib-mol
= 19.16
0.1
Number of
Transfer
Units = ln
(NOG)
J \
Hal — Cone
1^0.02 * Hal — Cone
* (i -I/AF) + (I/AF)
NOG =
5. Determine the height of a transfer unit.
The height of a transfer unit is the height of one
theoretical equilibrium stage. This number is determined from
the vapor and liquid flow rates per cross-sectional area of
the tower and from the Schmidt numbers
liquid.
8,9
Vapor Flowrate
per Cross-Sectional
Area
hr
= (0.3168
of the vapor and the
3600 sec
ft2-sec
hr
= 1140.48
Ib
ft2 • hr
B-22
-------
Liquid Flowrate
per Cross-Sectional
Area
ft2 • hr
(102.23 lb"m°M */18 —^-)
_\ hr / \ Ibmol;
0.842 ft2
= 2185.4
lb
ft2 • hr
Height of a
Transfer Unit
/ 1
+ ( AF
HT
* HL
where:
HG = b * (Vapor Flowrate per Area)C ^
(Liquid Flowrate per Area)a I Number /
HL = Y * ( Liquid Flowrate per Areaf ,
\ Viscosity of the Liquid / Number )
' 5
The values b, c, d, Y, and s and the Schmidt numbers are all
constants which are dependent on the packing and the liquid
and vapor flow rates through the tower. The height (HOG) of a
transfer unit is calculated below for the example facility.
3.82 * 1140.48
1b
HG =
°-41
* (0.809)
2185.4
lb
D
^
hr
-rr
. 45
= 1-93
HL = 0.0125 *
= 1.13
2185.4
lb
hr
0.85cp * 2.42
lb
hr • f t • cp t
0.22 0.5
* (381)
HOG = 1.93 +
* 1.13| = 1.989 ft
B-23
-------
6. Determine the height of the column.
The height of the column is determined from the number of
transfer units and the height of a transfer unit.
Height (ft) = (HOG * NOG) +2 + (0.25 * Diameter)
Height (ft) = (1.989 * 4.072) + 2 + (0.25 * 1.035)
= 10.36 ft
7. Determine the pressure drop across the tower.
The pressure drop is related to the height of the tower,
the liquid and vapor flow rates per cross-sectional area, and
the liquid and vapor densities.
Pressure _ NOG * HOG * g * (10~8)
Drop ^~2
* (10) (r * L'VpL) *
Pv
Where:
NOG = Number of Transfer Units
HOG = Height of a Transfer Unit (ft)
pL = Density of the liquid (lb/ft3)
pV = Density of the vapor (lb/ft3)
L" = Liquid Flow rate per Area (Ib/ft2«hr))
G" = Vapor Flow rate per Area (Ib/ft2»hr))
5.2 = Conversion between lb/ft2 and inches of H2O
The values of g (11.13) and r (0.00295) are constants
9
dependent on the tower design parameters.
B-24
-------
Pressure . 11.989^4.072| . 11-13 . (10-8) ,
0.00295 * 2185.4
(10)\ 62.2
(1140.48^ =3_87 inchesH9Q
0.0739
8. Calculate the weight of the tower.
As explained in Section 3.1.2.2 of the text, the weight
of the tower is dependent on the height and diameter of the
tower.
Weight = (48 * Diameter (ft) * Height (ft)) + 39 * (Diameter)2
Weight = (48 * 1.035 * 10.36) + 39 * (1.035)2 = 556.5 Ib
9. Calculate the capital cost of the scrubber.
The capital cost of the tower and associated equipment is
calculated from the design parameters. The procedure follows
the procedure presented in Section 3.1.2.2 of the text. All
costs are reported in July, 1989 dollars.
Tower Cost = (l.900604 * (wt (1*>M ' \ * $1000 * Cost Index
\ \1000 Ib/ / $298.2
0 93839
Tower Cost = (l.900604 * ( 556'5*b) 1 * $1000 * $355.9
\ \ 1000 Ib / / $298.2
= $1308.76
The volume of packing is needed to determine packing
costs.
The cost of the ducting and fans are based on 100 feet of
24-inch diameter ducting that is 1/8 inch thick carbon steel
and has two elbows per 100 feet.
B-25
-------
= * *
-------
EEC = $1308.76 + $80.03 + $3910.80 + $489.27 +
$2684.99 + $5000.00 = $13,473.85
The purchased equipment cost includes the EEC and
instrumentation, freight, and sales tax.
Instrumentation =0.1* BEG
Freight = 0.05 * BEG
Sales Tax = 0.03 * BEG
PEC = 1.18 * BEG = 1.18 * $13,473.85 = $15,899.14
The total capital cost of the scrubber system
incorporates an installation factor of 2.2.
TCI - 2.2 * PEC = 2.2 * $15,899.14 = $34,978.11
The total capital investment of the incinerator and
scrubber system is the sum of the two capital costs.
TCI = Incinerator TCI + Scrubber TCI
= $87,718.44 + $34,978.11 = $122,696.55
B.2 TOTAL ANNUAL COST
The total annual cost includes the direct and indirect
costs associated with the incinerator and scrubber. Direct
costs include costs such as labor, utilities, and maintenance,
and the annual cost associated with the nitrogen blanketing
system. Indirect costs include overhead, taxes, insurance,
administrative costs, capital recovery, and the annual costs
associated with maintaining the transfer rack and vehicles.
B.2.1 Incinerator Annual Cost
1. Calculate the annual incinerator direct cost.
The direct annual cost associated with the incinerator is
the sum of operating and supervising labor costs, maintenance
labor and material costs, and utility costs. Labor wages are
reflective of July, 1989.
B-27
-------
la. Calculate operating and supervisory labor costs.
It is assumed that the supervisory labor cost is 15% of
the operating labor cost.
Operating
+ Supv Labor = (1.15) * " V." * (Operating Hrs —\
Cost \8 hr shift/ \ * * yr/
* /Labor Wage -£-\
\ hr/
.
8 / \ yr/ \ hr
= $582.72/yr
Ib. Calculate maintenance material and maintenance labor
costs.
It is assumed that the cost of maintenance material is
equal to the cost of maintenance labor.
Maintenance
Maintenance Labor = $556.62/yr
Maintenance Material = $556.62/yr
ic. Calculate utility costs.
The utility costs associated with the incinerator include
the costs of natural gas and electricity.
Natural Gas
Cost
= / $3.03 \ / Natural Gas \ #
\1000 SCF/ \Flowrate SCFM/
so - * operating Hours — \
* a yr/
( $3'03 \ * (5.35 SCFM) * (60 ^\ * (614.2 **}
\ 1000 SCF/ \ hr / \ yr/
$597 .39/yr
The electricity costs are the costs associated with
running the incinerator fan.
B-28
-------
The pressure drop across an incinerator with zero heat
recovery is 4 inches water. The efficiency of the fan and
motor is assumed to be 60 percent.
(1.17 * 10~4 * Volumetric Flow (SCFM) *
Pressure Drop/Efficiency)
1.17 * 10~4 * 216.3 SCFM * —— \ = 0.1687 KW
0.6/
Electricity _ $0.0509 Pnwpr /™N * Operating /Hr\
Cost ~ KW . hr Fower * Hours [^j
= $°-0509 * 0.1687 KW * 614.2 — = $5.27/yr
KW • hr yr
The direct annual cost associated with the incinerator is
the sum of the costs presented above.
Incinerator
Direct Annual = Operating/Supv. Labor + Maintenance and
Cost Labor
Materials + Utilities
= $582.72 + $556.62 + $556.62 + $597.39 + $5.27
= $2298.62/yr
2. Calculate the indirect annual cost for the
incinerator.
The indirect annual cost associated with the incinerator
includes the overhead costs; taxes, insurance, and
administrative costs; and the capital recovery factor.
2a. Calculate the overhead cost.
The overhead cost is a function of the operating and
supervisory labor costs and the maintenance labor and material
costs.
Overhead = (Operating/Supv. Labor + Maintenance
Labor and Materials) * 0.6
Overhead = ($582.72 + $556.62 + $556.62) * 0.6 = $1017.58/yr
B-29
-------
2b. Calculate the taxes, insurance, and administrative
costs and the capital recovery. Capital recovery is based on
equipment life of 10 years and an interest rate of 10 percent.
The taxes, insurance, and administrative costs, and the
capital recovery costs are all dependent on the total capital
investment of the incinerator.
Taxes = 0.01 * TCI
Insurance = 0.01 * TCI
Administration = 0.02 * TCI
Capital Recovery Cost = 0.16275 * TCI
= °-20275 * TCI = 0.20275 * $87,718.44
= $17,784.91/yr
The indirect annual cost for the incinerator is the sum
of the above costs.
Incinerator
Indirect Annual = Overhead + Taxes + Insurance +
Cost
Administration + Capital Recovery
= $1017.58 + $17,784.91
= $18,802.49/yr
The total direct annual cost of the incinerator is the
sum of the direct annual cost and the indirect annual cost.
Incinerator
Total Annual = Direct Annual Cost + Indirect Annual Cost
Cost
= $2298.62 + $18,802.49 = $21,101 .11/yr
B.2.2 Scrubber Annual Cost
1. Calculate the direct annual cost.
The direct annual cost associated with the scrubber
includes the operating and supervisory labor costs, the
maintenance material and labor costs, and the utility costs.
It is assumed that the scrubber operates for the same number
B-30
-------
of hours per year as the associated incinerator. Labor rates
are reported in July, 1989 dollars.
la. Calculate the operating and supervisory labor.
It is assumed that the supervisory labor cost is
15 percent of the operating labor cost.
Operating/
supervising - ,1.15) .
(hr/yr) Wage \hr/
- Iil5 * 0-5 hrs t 614_2 hr „ $13.20
8 hr yr hr
= $582.72/yr
Ib. Calculate the maintenance material and maintenance
labor cost.
It is assumed that maintenance material is equal to
maintenance labor.
Maintenance - • 5 ?" ) * (°PS05rsng ^) * --
8 hr shift/ \ Hours yr / \ Wage yr
*/614'.2
.
8 hr / \ yr/ \ hr
Maintenance Labor = $556.62/yr
Maintenance Material = $556.62/yr
Ic. Calculate utility costs.
The scrubber utility costs include water and electricity
costs. The electricity costs depend on the vapor flow and
pressure drop through the column.
B-31
-------
Water
Cost
water (Ib/hr)
8.34 (Ib/gal)
/1QO 03 Ibmol
. \ hr
ft 74 -
t $0.22 ^
1000 gal
+19 lb \
18 Ibmol)
lb
• 614 2 hr
yr
t $0.22 .
1000 gal
t 614 -> hr
yr
gal
= $29 . 81/yr
Vapor
EleCrniCity - °-°°02 * Thogh * Prg^re (inches H20) *
crubbe uro
(SCFM)
rn
cost Scrubber
Operating / hr \ ^ $0.0509
Hours \yr) KW • hr
= 0.0002 * 216.3 SCFM * 3 . 87 inches *
614.2 M *10^9
yr KW • hr
= $5.23/yr
The direct annual cost of the scrubber is the sum of the costs
presented above.
Scrubber
Direct Annual = Operating Supv. Labor + Maintenance Labor and
Cost
Materials + Utilities
= $582.72 + $556.62 + $556.62 + $29.81 + $5.23
= $1731.00/yr
2. Calculate the scrubber indirect annual cost.
The scrubber indirect annual costs include the overhead,
taxes, insurance, administrative costs, and capital recovery.
2a. Calculate the overhead cost.
The overhead cost is a function of the operating and
supervisory labor costs, and the maintenance labor and
materials costs.
B-32
-------
Overhead = (Operating/Supv. Labor + Maintenance Labor and
Materials) * 0 . 6
= *$582.72 + $556.62 +• $556.62) * 0.6
= $1017.58/yr
2b. Calculate the taxes, insurance, and administrative
costs, and the capital recovery costs. These costs are
dependent on the total capital cost of the scrubber. Capital
recovery is based on equipment life of 10 years and an
interest rate of 10 percent.
Taxes = 0. 01 * TCI
Insurance = 0.01 * TCI
Administration = 0.02 * TCI
Capital Recovery Cost = 0.16275 * TCI
= 0.20275 * TCI = 0.20275 * $34,978.11
= $7091.81/yr
The indirect annual cost is the sum of the above costs.
Scrubber
Indirect = Overnead + Taxes + Insurance +
Cost
Administration + Capital Recovery
= $1017.58 + $7121.42 = $8139.00/yr
The total annual cost of the scrubber is equal to the sura of
the direct annual cost and the indirect annual cost.
Scrubber
;Tr2^\ = $1731.00 + $8139.00 = $9870. 00/yr
Cost
B-33
-------
B.2.3 Total Annual Cost
The total annual cost includes the total annual cost of
the incinerator and scrubber, the annual costs associated with
the nitrogen blanketing system, and the annual costs
associated with maintaining the transfer rack and vehicles.
1. Calculate the annual cost associated with the nitrogen
blanketing system.
This cost is dependent on the total material throughput
to the transfer vehicle.
Tank Truck
Nitrogen = $274 * Vehicle Annual Throughput (MMgal/yr)
Annual Cost
= $274 * 12.53 = $3433.22/yr
2. Calculate the annual cost of maintaining the vapor
collection system (rack annual cost) . The cost ($200) is
adjusted to July, 1989 dollars using CE Construction Labor
indices.
Annual = $200 * 26-^ - $244.18/yr
Cost 22°
3 . Calculate the vehicle annual cost for vapor tightness
testing.
The vehicle annual cost is dependent on the vehicle fleet
size, which is dependent on the number of loads. The number
of loads is equal to the total maximum production capacity of
all transferred chemicals divided by the vehicle capacity.
The tank truck and tank car vehicle capacities are equal to
10,000 gallons and 20,000 gallons, respectively.
Number _ Maximum Capacity (gal/yr)
of Loads vehicle Capacity (gal)
Number of 15-26 * 106
Tank Truck = - i±- = 1526 loads
Loads 10,000 gal
If the number of loads is greater than 1463 and less than or
equal to 1646, then the number of vehicles is equal to nine.
B-34
-------
The cost ($300) is also adjusted to July, 1989 dollars using
CE Construction Labor indices.
Vehicle .,„„ „,„ .. Vehicle
Annual = 30° * 268'6 * Fleet
Cost 22° Size
Tank Truck -nn 9f.ft ,
Vehicle = 30° * 268'6 * 9 = $3296.45/yr
Annual Cost 22°
The total annual cost is the sum of the incinerator annual
cost, the scrubber annual cost, the nitrogen annual cost, and
the annual cost of the transfer rack and vehicles.
Total
Annual = $21,101.11 + $9870.00 + $3433.22 +
Cost
$244.18 + $3296.45
= $37,944.96/yr
Thus, the total annual cost of controlling the tank truck rack
at the example facility is approximately $37,900/yr.
B-35
-------
REFERENCES
1. Compilation of Air Pollutant Emission Factors, Volume 1:
Stationary Point and Areas Sources. U.S. Environmental
Protection Agency, Office of Air Quality Planning and
Standards. Research Triangle Park, NC. Publication No.
EPA/AP-42. September 1985. pp. 4.4-1 through 4.4-17.
2. U.S. Environmental Protection Agency, Office of Air
Quality Planning and Standards. OAQPS Control Cost
Manual. Fourth Edition. Section 3.9. EPA-450/3-90-006.
Research Triangle Park, NC. January 1990.
3. Memorandum from Pandullo, R.F., Radian Corporation, to
Barbour, W., and D. Stone, Radian Corporation; L. Evans,
and B. Rosensteel, EPA/CPB; and B. Vatavuk, EPA/SDB.
April 27, 1990. Summary of April 11 meeting to discuss
thermal incinerator cost issues.
4. U.S. Environmental Protection Agency, Air and Energy
Engineering Research Laboratory. Handbook—Control
Technologies for Hazardous Air Pollutants.
EPA-625/6-86-014. Research Triangle Park, North
Carolina. September 1986.
5. Memorandum from Ferrero, B. , Radian Corporation, to HON
project file. February 5, 1992. Estimating liguid to
vapor flow rate ratios in scrubber columns.
6. Treybal, R.E. Mass-Transfer Operations, Third Edition.
New York, McGraw-Hill Classic Book Company. 1980.
p. 196.
7. Memorandum from Ferrero, B., Radian Corporation, to HON
project file. February 5, 1992. Development of the
slope of the equilibrium line and the absorption factor
for acid gas scrubber design.
8. Memorandum from Barbour, W., Radian Corporation, to HON
project file. April 20, 1990. Estimating liquid and
vapor Schmidt numbers for acid streams.
9. Danielson, John, A. Air Pollution Engineering Manual,
Second Edition. Air Pollution Control District County of
LA. U.S. Environmental Protection Agency, Office of Air
Quality Planning and Standards. Research Triangle Park,
NC. May 1973. p. 214-217.
10. Memorandum from Self, P., Radian Corporation, to HON
project file. February 5, 1992. Estimating scrubber
stack cost.
11. Memorandum from Scott, K., Radian Corporation, to HON
Project file. February 3, 1992. Estimating vehicle
fleet size for HON transfer.
B-36
-------
APPENDIX C
EXAMPLE COSTS FOR INSTALLATION OF A
REFRIGERATED CONDENSER TO A
FIXED ROOF TANK
This is an example calculation to determine the total
annual costs associated with controlling an example storage
tank with a refrigerated condenser having a removal efficiency
of 95 percent. The purpose of this appendix is to demonstrate
the approach used in the HON analysis. In the calculations
below, all significant figures have been retained until the
final calculation to make it easier for the reader to follow
the calculation and to avoid potential error due to round off
of intermediate calculations. It should not be inferred that
the intermediate results represent the actual number of
significant figures.
The example tank has a capacity of 20,000 gallons and
stores vinylidene chloride (CH2=CCl2). Design and cost
assumptions are presented in Tables C-l and C-2.
1. Determine the filling rate for the example tank which
is dependent upon the tank capacity. Table C-3 presents tank
filling rates which were developed from engineering judgement
for five tank capacity ranges. The example model tank has a
storage capacity of 20,000 gallons; therefore, its filling
rate is 500 gallons/minute.
2. Calculate the annual number of hours a tank
experiences working losses. Working losses occur during tank
filling operations when the incoming fluid displaces vapors in
the tank head space. Therefore, working hours are the number
of hours spent each year filling a tank.
C-l
-------
TABLE C-l. DESIGN AND COSTa ASSUMPTIONS
Parameter Description
(Framework Variable Names)
Value
Annual tank breathing hours
(b hours)
Storage tank operates at
standard conditions
Volume of ideal gas at
standard conditions
Molecular weight of air
Condenser HAP removal
efficiency
(rem eff)
Condenser mechanical work
efficiency
Electricity costs
(eleccost)
Water costs
(watercst)
Labor costs
(laborcst)
Maintenance labor costs
(mtlbrcst)
8,760 hr/yr
Total pressure of 760 mmHg
and a temperature of 25 °C
392 scf/lb-mol of gas
29 Ib/lb mol
0.95
0.85
$0.0509/kW-hr
$0.00022/gal
$13.20/hr-labor
$14.50/hr-labor
aCosts are given in July 1989 dollars.
C-2
-------
TABLE C-2. MODEL TANK DESIGN PARAMETERS
Parameter Description
Units
Value
Number of tanks
Tank volume (tank_size)
Annual tank throughput
(vol_per_tk)
Tank orientation and type
Tank diameter (tank_dia)
Tank height
Average tank vapor space
height
Adjustment factor for small
diameter tanks (C
-------
TABLE C-3. TANK FILLING RATES
Tank Capacity Filling Rate
(gallons) (gallons/minute)
Capacity < 10,000 250
10,000 < Capacity < 20,000 500
20,000 < Capacity < 40,000 750
40,000 < Capacity < 200,000 1,000
Capacity > 200,000 2,000
C-4
-------
Workina Hours - annual volume throughput of a single tank
filling rate
w hours = vol— Per— tk (gal/yr)/ 1 hr \
w — iiLJvai. o — _ 7- •-_ _ ----- / ~ , : ^M ~ ~ : I
fill rate (gal/min) \60 mm/
w_hours = 2,472,727 (gal/yr), 0_0167 Jir_
500 (gal/min) mm
w hours = 82.589 hr/yr
3. Calculate the number of tank turnovers per year,
Turnovers = annual throughput of a single tank
tank capacity
turn = vo1—Per—tk (gal/yr)
tank—size (gal)
turn = 2,472,727 (gal/yr)
20,000 (gal)
turn = 123.64 turnovers/yr
4. Calculate uncontrolled working and breathing losses
from the example tank using AP-42 emission equations adjusted
to yield results in Mg/yr.
4a. Fixed roof tank working losses:
Fixed Roof HAP HAP
Tank . nfiQ „ nn-e * Molecular Vapor Pressure
Working ~ •>-'"** x 10 * Weight * at 25°C
Losses (Ib/lb-mole) (psia)
Annual
Volume Tank
* Throughput * Turnover *
for a tank Factor
(gal/yr)
C-5
-------
work — u = 1.089 X 10 8 * mol— weight * vp_psia
* vol_per_tank * / f ° * turn] * i.o
\ (6 * turn) /
work— u = 1.089 X 10~8 * 96.94 * 11.6279
* 2,472,727 * (180 +123. 64) \ ,
\ (6 * 123.64) /
work _ u = 12.424 Mg/yr
4b. Fixed roof tank breathing looses:
Fixed Roof HAP Partial
Tank _ -5 Molecular Pressure Ratio
Breathing - 1-025 x 10 * Weight * "f SJIU- F® R^ t:10
Losses (Ib/lb-mole) of ^P to Air
Tank Average Tank Average Ambient
* Diameter * Vapor Space *
(ft) Height (ft)
Tank
Tank Adjustment
* Paint * Factor for *
Factor Small Diameter
Tanks
breath— u = 1.025 X 10~5 * mol— weight
+ / _ vp— psia _ \0.68
\total pressure - vp — psia/
* (tank— dia)1-73 * (dh)°-51 * (dt)°-50
* (Fp) * (Cd) * (Kc)
C-6
-------
breath—u = 1.025 X 10~5 * 96.94
11.6279 \0.68
14.7 - 11.6279
* (15)1'73 * (7.5)0'51 * (20)°'5°
* (1.3) * (0.7306) * (1.0)
breath _ u = 3 . 158 Mg/yr
5. Determine the mass flow rate of HAP emissions into
the condenser based on working and breathing losses from the
tank.
rT ^t-^^i ^ ^ Uncontrolled
wSkiSg'LosIes - ""icing ix, .. Emissions
Entering Condenser Working Hours
work in = work-u (Mg/yr) /106g\/ 1 lb \
w_hours (hr/yr) I Mg A453.6g)
work—in = W0lk— ^^ (Mg/yr) * 2204.6 (Ib/Mg)
w — hours (hr/yr)
work— in = 12.424 (Mg/yr) * 2204. 6 (Ib/Mg)
82.589 (hr/yr)
work in = 331.642 Ib/hr
n, j uncontrolled
Uncontrolled breathing loss emissions
breathing emissions = :" ^—. r
entering condenser breathing hours
_ breath—u (Mg/yr) * 2204.6 (Ib/Mg)
b—hours (hr/yr)
. .. . _ 3.158 (Mg/yr) * 2204.6 (Ib/Mg)
bth-in 8760 (hr/yr)
bth in = 0.7948 Ib/hr
C-7
-------
6. Calculate the HAP concentration in working losses at
the inlet to the refrigerated condenser.
rg of ">-n«>les of
losses at the - HAP in working losses / 106 ppm\
condenser inlet total number of \ 1 /
(concen _ w) Ib-moles in working
loss gas stream
6a. First, calculate the number of Ib-moles of HAP in
the working losses:
HA^workfng = work-u (Mg/yr) * 2204.6 (Ib/Mg)
losses m°l—weight (Ib/lb-mol)
6b. Then, calculate the total number of Ib-moles in the
working losses:
Total Ib-moles _ vol—per—tk (gal/yr)/0.1337 ft3
in working losses 392 (scf/lb-mol) \ gal
Total Ib-moles „. _„ <-u * o ,,-11 -i n-4/lt> mol\
in working losses = vol-per_tk * 3.411 x 10 4
6c. Divide equation 6a by equation 6b:
/work—u * 2204. 6\ + /106 ppm\
concen-w = \ mol-weight / \ 1 /
00 CS vol-per_tk * 3.411 X ID'4
concen_w =
f
(mol — ^weight) (vol — per — tk)
C-8
-------
concen-w = . * (12.424)
(96.94) *(2,472,727)
concen w = 334,823 ppmv
7. Calculate the HAP concentration in the breathing
losses at the inlet to the refrigerated condenser.
a number of Ib-moles of
lossJS at the = HAP in breathing losses / 10* ppmv
condenser inlet ,, total number of \ 1 /
(concen _ b) Ib-moles in breathing
loss gas stream
7a. First, calculate the number of Ib-moles of HAP in
the breathing losses:
Ib-moles of HAP = breath—u (Mg/yr) * 2204.6 (Ib/Mg)
in breathing losses mol—weight (Ib/lb-mol)
7b. Then, calculate the total number of Ib-moles in the
breathing losses. It was assumed that the HAP liquid phase
was in equilibrium with the HAP gaseous phase and Raoult's Law
applies. Thus, the ratio of the HAP partial pressure divided
by the total pressure is equal to the HAP mole fraction in the
gaseous phase. Therefore, total pressure (760 mm Hg) and
vap pr 25C (HAP partial pressure in mmHg) are shown below in
units of Ib-moles.
Total
Ib-moles = breath—u (Mg/yr) * Total Pressure (Ib-mol)
breathing mol—weight (Ib/lb-mol) * vap—pr—25C (Ib-mol)
losses
* 2204. 6 O*\
\Mg/
C-9
-------
7c. Divide equation 7a by equation 7b:
rbreath—u * 2204.6 (Ib/Mg)
concen—b = I mol—weight |
rbreath—u * total pressure *2204.6 (Ib/Mg)
[ mol—weight * vap—pr—25C
106 ppm\
1
concen b = 1316 * vap pr 25C
concen b = 1316 * (601.17)
concen b = 791,140 ppm
8. Calculate total volumetric flow rates at standard
conditions (760 mmHg and 25°C) for working losses and
breathing losses at the refrigerated condenser inlet.
Volumetric flow
rate of working =
losses storage tank
w flow = fillrate (gal/min) * 0.1337 (ft3/gal)
w flow = 500 (gal/min) * 0.1337 (ft3/gal)
w flow = 66.85 ft3/min @ 25°C
Uncontrolled HAP emissions
Volumetric flow from breathing losses at
rate of breathing _ the condenser inlet
losses at the ~ HAP HAP concentration
condenser inlet molecular * in breathing losses
weight at the condenser inlet
b_fiow = / bth—in (Ib/hr) \ + / 106 (Ib-mol total)
,mol—weight (Ib/lb-mol) / \concen-b (Ib-mol HAP)
392 scf\ ^ I 1 hr \
+ (392 scf\ „, / Ihr \
\ Ib-mol / \60 min/
C-10
-------
b_flow = (6,533,333 scfm) (bth—in)
(mol—weight) (concen—b)
b_flow = (6,533,333) (0.7948)
(96.94) (791,140)
b flow = 0.06771 ft3/min @ 25°C
9. Calculate the mass flow rate of carrier gas (e.g.,
air) which will enter the refrigerated condenser along with
the working and breathing losses.
Carrier gas
mass flow rate Volumetric flow Concentration
for working = rate of working * of air in
losses at the losses working losses
condenser inlet
lp6 " °°cen-w lb-mol air U
lb-mol total /J
w_air = w_flow (scfm) * " -w (
\
1 Ib-mol total \ / 29 Ib air
392 scf
\ ^ / 29 Ib air \ ^ /60 min\
/ (ib-mol air/ \ 1 hr /
w—air = (4.439 * 10"6lb/hr) (w_flow) (106 - concen—w)
w—air = (4.439 * 10~6 Ib/hr) (66.85) (106 - 334,823)
w air = 197.39 Ib/hr
Carrier gas
mass flow rate Volumetric flow Concentration
for breathing = rate of * of air in
losses at the breathing losses breathing losses
condenser inlet
C-ll
-------
b_air = b_flow (scfm) * [ ^6 - concen-b / Ib-mol air vi
[ 10*\ Ib-mol totaljj
„, /1 Ib-mol total\ + / 29 Ib air \ + / 60 min\
\ 392 scf / \lb-mol air) \ l hr )
b—air = (4.439X10'6 (lb/hr)) (b—flow) (106 - cone en—b)
b—air = (4.439X10"6 (lb/hr)) (0.06771) (106 -791,140)
b air = 0.06278 lb/hr
10. Calculate the mass flow rate of HAP working and
breathing losses removed by a refrigerated condenser with a
removal efficiency of 95%.
Working loss Working loss
HAP emissions _ HAP emissions + condenser
removed by entering the efficiency
the condenser condenser
work rem = work in (lb/hr) * rem eff
work rem = 331.642 (lb/hr) * 0.95
work rem = 315.060 (lb/hr)
Breathing loss Breathing loss
HAP emissions _ HAP emissions + condenser
removed by entering the efficiency
the condenser condenser
bth rem = bth in (lb/hr) * rem eff
bth rem = 0.7948 (lb/hr) * 0.95
bth rem = 0.7551 lb/hr
C-12
-------
11. Calculate the mass flow rate of HAP working and
breathing losses at the condenser outlet.
Working loss Working loss Working loss
HAP emissions _ HAP emissions _ HAP emissions
exiting the entering the removed by the
condenser condenser condenser
work out = work in (Ib/hr) - work rem (Ib/hr)
work out = 331.642(lb/hr) - 315.060 (Ib/hr)
work out = 16.582 Ib/hr
Breathing loss Breathing loss Breathing loss
HAP emissions _ HAP emissions _ HAP emissions
exiting the entering the removed by the
condenser condenser condenser
bth out = bth in (Ib/hr) - bth rem (Ib/hr)
bth out = 0.7948 (Ib/hr) - 0.7551 (Ib/hr)
bth out = 0.0397 Ib/hr
12. Calculate the vapor pressure of the non-condensed
working losses exiting the condenser. The HAP vapor pressure
will be the partial pressure as determined from Raoult's Law.
[Note: The cost of a refrigerated condenser system is a
function of the temperature to which the HAP vapor must be
cooled in order to condense. Systems designed to achieve
lower temperatures are more expensive because they can require
different coolants and materials of construction. Also,
different operating temperatures can increase the amount of
energy that must be put into the system.
For a given HAP removal efficiency, the condenser outlet
temperature is a function of the level of saturation of the
stream entering the condenser and, therefore, the vapor
pressure of the stream exiting the condenser. Saturated
streams such as tank breathing losses condense more easily;
lower temperatures are required to condense unsaturated
C-13
-------
streams such as tank working losses. As a conservative
estimate of condenser outlet temperature (and therefore,
system cost) , the condenser outlet temperature will be
calculated based on the working losses only. ]
losses .
at. une
condenser outlet
vpout = 760 mm Hg * ( — lb-mol HAP v
\lb-mol air + lb-mol HAP/
/ work—out (Ib/hr) \
vpout=760mmHg* 1mol-weight (Ib/lb-mol) )
w— air (Ib/hr) \ +
2 9 lb/ lb-mol
\
/
— out-r —
work — out
r —
hr/
mol— weight/— -r - =-)
^ \ lb-mol I)
vpout = 760 ram Hg *
16.582 Ib/hr \
96.94 Ib/lb-mol/
197.29 Ib/hrV + / 16.582 Ib/hr \
29 Ib/lb-mol) \96.94 Ib/lb-mol/
vpout = 18.640 mm Hg
13. Calculate the temperature at the outlet of the
refrigerated condenser (tout c) using Antoine's equation:
Iog10 (vpout, mm Hg) = A - (tout_GBoc) +Q
Where: A, B, and C are Antoine coefficients.
It is assumed that the gaseous (noncondensed) HAP
emissions exiting the condenser were in equilibrium with the
condensed HAP emissions. Thus, the temperature calculated
C-14
-------
from Antoine's equation will be the temperature at the
condenser outlet and the condensation temperature necessary to
yield a 95% removal efficiency.
Condenser outlet _ Antoine's equation solved
temperature for temperature
T3
(vpout) = A -
(tout— c) + C
Iog10 (vpout) - A =
tout—c + C = - B
Iog10 (vpout) - A
tout—c = -C - B
(vpout) - A
tout-c = - (antcof_c) - ' antcof_b \
log-LQ (vpout) - antcof—a/
tOUt-C = - (237.20) -' 1099.40
Iog10 (18.640) - 6.972
tout C = -44.38°C
14. Convert the condenser outlet temperature from
degrees Celsius to degrees Kelvin. The condenser outlet
temperature is assumed to be the same for both working and
breathing losses.
C-15
-------
Condenser outlet temperature ^ ., ., .
in H^crrpp=! K^i-Jin Condenser outlet
^r^rlfing^r111 ' ^-^1^ +2?3'15
breathing losses aegrees celcius
temp—w—k
or = tout—c + 273.15 (°K)
temp—b—k
temp—w—k
or = (-44.38 + 273.15) (°K)
temp—b—k
temp—w—k
or = 228.77°K
temp—b—k
15. Convert the condenser outlet temperature from
degrees Celsius to degrees Fahrenheit.
Condenser outlet Condenser outlet
temperature in = temperature in
degrees Fahrenheit degrees Celcius
tout f = (tout c * 9/5) + 32
tout f = (-44.38 * 9/5) + 32
tout f = -47.88 °F
16. Calculate the contribution to the condenser heat
load made by the HAP sensible heat loss from working and
breathing losses:
C-16
-------
HAP sensible rate of HAP Temperature
heat loss _ HAP working . molar . iT^,,®*®1^®
from - losses at * heat * ^^ween ^? t-
working losses the condenser capacity condenser inlet
inlet and Outlet
w_ht_gas = ( ~- ht_cap
mol—weight lb/lb-mol/ \ Ib-mol • °F
(25°C - tout—c) * [l.8 4^]
,, v,*. / 331.642 lb/hr \ /00
w—ht—gas = ^ ^ /11_/ r- 23
\96.94 lb/lb-molj \
Btu
Ib-mol
* (25°C - [-44.38°C]) * (l. 8 45)
w ht gas = 9827 Btu/hr
HAP sensible Molar flow rate
heat loss from of HAP breathing .
breathing = losses at the * catv COnden inet
ir-icjccic <-1c»i- cd.pa.cicy coriuensei iiiiec
losses condenser inlet and outlet
b-ht_gas = , ~ - ht-cap
, - -1u n 0
mol-weight lb/lb-mol/ \ Ib-mol - °F
(25°C - tout— C) l-84
C-17
-------
b_ht
eras = / 0-7948 Ib/hr \ /
~gaS (96.94 Ib/lb-mol) (23
Btu
lb-mol-°F
(25°C - [44.38 °C]) * [l. 8 -^
b _ ht _ gas = 23.55 Btu/hr
17. Calculate the contribution to the condenser heat
load made by the air sensible heat loss from working and
breathing losses.
Air sensible
heat loss
from working
losses
Molar flow
rate of
air in working
losses at the
condenser
inlet
,. mnlar
t
capacity
Temperature
difference
between tne
condenser
inlet and
outlet
w—ht—air - / w—air Ib/hr \ [6_713 (299.15 - temp—w_k)
\29 Ib/lb-mol/
4699*10~4, o o,
+ a-b^ ±H— 298.152 - temp—w—k2
o
4.994 * 10
-10
(298-15
4 _
\gmol /
* /3.9685 * 10~3
cal
453>6
\
Ib-mol
c-18
-------
w— ht— air = 0.062 * (197.39) *
[6 .713 (298.15 - 228.77)
+ 4.699 * 10"4(298.152 _ 228.772)
+ 1'147 * 10 - (298. 153 - 228.77s)
- 4.994 * 10-10(298.154 . 228.77^)]
4
w _ ht _ air = 5865 Btu/hr
Air sensible Molar flow rate
heat loss of air in Air molar
from = breathing losses * heat *
breathing at the condenser capacity
outlet
b— ht— air = / b— ^ ir lb/hr \ [6-713 (298.15 _ temp— b— k)
\ 29 lb/lb-mol/
+ 4.699 * 10"4(298.152 _ temp-b-k2)
£
* 10"6(298.153 _
-10
- 4.994 * 10"-LU(298.154 . temp-b-k*)] (-2*1 \
4 \gmol /
\
* 3.9685 * 10~J^^ * 453.6
.
cal/ \ lb-mol/
C-19
-------
b— ht— air = 0.062 * (0.06278) *
[6.713 (298.15 - 228.77)
+ 4.699 * 10-4(298.152 -228.772)
+ 1-147 * 10 - (298. 153 -22S.773)
- 4.994 * 10-10(298.154 -228.774)]
4
b _ ht _ air = 1.865 Btu/hr
18. Calculate the contribution to the condenser heat
load made by the HAP heat of vaporization.
HAP heat of
vaporization _ kna * mp heat of
from working ' ^Ses frSm vaporization
losses the condenser
w—ht—lq = [work—rem —— | (ht—vap
\ hr/ \ Ib
w—ht_lq =
..
\ hr/ \ Ib
w—ht—lq = 40,851-5^
HAP heat of Mass flow rate
vaporization _ of condensed HAP ^ HAP heat of
from breathing breathing losses vaporization
losses from the condenser
C-20
-------
b— ht— Iq = (bth— rem -^-\ (ht— vap
\ hr/ \
b— ht— Iq = /0. 7551-^/129. 660
\ hr/\
Ib
Tr
Ib
b—ht—Iq = 97.9
hr
19. Calculate the total heat load that the condenser
must remove from the entering working and breathing losses at
standard conditions (25°C and 760 mm Hg) to achieve a
condenser exit temperature of -44.38°C.
HAP sensible Air sensible
heat loss from heat loss from
Total the temperature the temperature Heat of
heat = difference + difference + vaporization
load between the between the of the HAP
condenser inlet condenser inlet
and outlet and outlet
Total condenser heat load from working losses:
w ht tot = w ht gas + w ht air + w ht Iq
w ht tot = 9827 Btu/hr + 5865 Btu/hr + 40,851 Btu/hr
w ht tot = 56,543 Btu/hr
Total condenser heat load from breathing losses:
b ht tot = b ht gas + b ht air + b ht Iq
b ht tot = 23.55 Btu/hr + 1.865 Btu/hr + 97.91 Btu/hr
b ht tot = 123.33 Btu/hr
C-21
-------
20. Calculate the refrigeration capacity required from
the condenser total heat load. Refrigeration is defined as
the quantity of energy to melt one ton of ice at 32°F in 24
hours, or 12,000 Btu/hr-ton.
Refrigeration
capacity required
for the condenser
heat load due to
working losses
Total condenser
12,000
'
Btu
hr-ton
w-refrig =
12,000
BtU
hr-ton
56,543
12,000
BtU
hr-ton
w _ refrig =4.71 tons
b_refrig =
b_ht_tot
12,000
BtU
hr-ton
123.33
= J
12,000
BtU
hr-ton
b _ refrig = 0.0103 tons
21. Calculate the number of hours the condenser
processes working losses assuming that the condenser operates
1/2 hour prior and 1/2 hour after each tank filling as well as
during tank filling.
Condenser Hour s per year
run time for = spent filling +
working losses a tank
ad
opeaton per
° filling
C-22
-------
w_run = w-hours M\ + (i ^r \ (turn tuinovers\
yr/ \ turnover/ \ yr /
^1 + fl turnovers
w_run = [82.589 + l - - 123.64
yr/ \ turnover/ \ yr
w—run = 206 .23 —
yr
NOTE: If w run >_ 8760 hr/yr then assume w run = 8760
hr/yr.
22. Calculate the number of hours the condenser operates
to process breathing losses. It is assumed that when the
condenser is not processing working losses, it operates
continuously to process breathing losses.
Condenser Number of hours condenser run
run time tank experiences _ ; .~~~~%n^
for breathing ~ breathing workinS losses
losses losses working losses
b—run = /b—hours —1 - /w— run^-
(b—hours — \ - /v
I yr/ I
b—run = (8760 — \ - /206.23 —\
yr/
b—run = 8553.77 —
yr
23. Calculate the basic equipment costs (EEC) for a
refrigerated condenser system based on the maximum operating
limit of the system, which occurs during the control of
working losses. Table C-5 contains equations for calculating
C-23
-------
EEC as a function of the condenser outlet temperature and
refrigeration capacity requirement.
For the example storage tank, the designed condenser has
an outlet temperature of -47.85°F (-44.36 °C); therefore, the
basic equipment costs are calculated from cost algorithm #12.
(
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C-25
-------
TABLE C-6. ESTIMATION OF TOTAL CAPITAL INVESTMENT FOR
A REFRIGERATED CONDENSER SYSTEM3'b
Component
Cost Component Cost Factor Costa
Direct Equipment Costs*3
Base Equipment Cost (BEG) TABLE 5 $89,461
Instrumentation Included in BEG
Sales Tax and Freight 8% of EEC $7,157
PURCHASED EQUIPMENT COST (PEC)
Installation Cost
Packaged System 15% of PEC
TOTAL INSTALLATION COST
TOTAL CAPITAL INVESTMENT (TCI)
$96,618
$14,493
$14,493
$111,111
aJuly 1989 dollars
kfiased on a HAP condensation temperature of -44°F and a
refrigeration requirement of 4.71 tons.
C-26
-------
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C-27
-------
Annual condenser Refrigeration Condenser
electricity costs _ requirement run time for
for controlling for working working
working losses losses losses
* elecTrVcity + Mechanical efficiency
costs of the Condenser
(eleccost
$
w—elec = (5 kw ^/w~refrig^/-- hr\J.\ kw-hr
ton/ \ icon; / yr 0.85
„ _n__ _ (5.0 kW/ton) (4.71 (ton)) (206.22 hr/yr) (0.0509 $/kW-hr)
6 6C " 0.85
w elec = $290.82/yr
NOTE: If w refrig had been less than the minimum value
(1.58 tons) listed in Table C-7, then setting w refrig equal
to the listed minimum value would yield a conservative
estimate of w elec.
Refrigeration Condenser
™i« requirement . run time for
for controlling = for breathing * breathing
losses losses
Mechanical
"l?^
condenser
r . eleccost
ton/ on yr
0 . 85
C-28
-------
As calculated in Step 20, b refrig - 0.0103 tons.
Because the lower bound of applicability for equation #4 in
Table C-7 is 1.58 tons, b elec must be calculated using this
minimum value. Since the condenser will actually be operated
at a much lower level (0.0103 tons), use of 1.58 tons gives a
conservative estimate of the electricity cost for the system.
b—elec = 15
.0 J2L\ (1.58 (ton)) (8553.78 **\ /$0.0509\
ton/ _l yr/ \ kW-hr /
0.85
b elec = $4,047/yr
26. Calculate the condenser's annual consumption of
electricity for controlling working and breathing losses.
Condenser's annual Annual condenser electricity
electofciifor = costs for controlling working losses
6 controlling0 Unit electricity costs
working losses
w—elec --)
w—kW_hr =
eleccost
290.82
w—kW—hr =
kW-hr
0.0509
w—kW—hr = 5,714
kW-hr
kW-hr
yr
Condenser's annual Annual condenser electricity
consumption of costs for controlling breathing losses
electricity for = TT ..—: —r^-r- ——
controlling Unit electricity costs
breathing losses
C-29
-------
b—elec I——
b—kW—hr =
eleccost
4047
b—kW—hr =
kW-hr
0.0509
-hr
b—kW—hr = 79,509 kw hl
yr
Condenser 'B total
annual consumption = contli wking and
of electricity breathing losses
tot__kW_hr = w kW hr + b kW hr
tot kW hr = (5,714 + 79,509) $/yr
tot kW hr = 85,223 $/yr
27. Calculate emission reduction as the working and
breathing losses removed by the refrigerated condenser.
emis—red =
work—rem w—hours bth—rem 8760-w—hours
/lb\ * /hrs\ + /lb\ * /hrs\
Ihr I yr / JhrJ ( yr
2204.6 —
Mg
emis
r~A - (315.060) * (82.589) + (0.7551 * [8760 - 82.589])
—]_ GO. — —• '"— — • ' - "•"" "• ~
emis red = 14.775 Mg/yr
C-30
-------
28. Calculate recovery credit as the product of emission
reduction and average market price.
rec credit = emis red (Mg/yr) * (1.57 $/kg) * (1000 kg/Mg)
rec credit = 14.775 * 1.57 * 1000
rec credit = 23,197 $/yr
29. Calculate total annual cost (TAG) using the approach
summarized in Table C-8. After rounding, the total annual
cost for controlling the example tank is approximately
$38,900/yr.
C-31
-------
TABLE C-8. ESTIMATION OF TOTAL ANNUAL COST FOR
A REFRIGERATED CONDENSER SYSTEM
Annual Annual
Cost Component Cost Factor Consumption Cost3
Direct Annual Costs
Utilities
Electricity $0.0509/kW-hr 85,223 kW-hr $4,338
Labor
Operating Laborb $13.20/hr 547.5 hrs $7,227
Supervision & 15% of Operating $1,084
Administration Labor
Maintenance
Laborb
Materials
TOTAL DIRECT
$14.50/hr
100% Of
Maintenance
Labor
ANNUAL COST (TDAC)
547.5 hrs $7,939
$7,939
$28,527
Indirect Annual Costs
Overhead 60% of all labor $14,513
and materials
General &
Administrative
Property Taxes 1% of TCI $1,112
Insurance 1% of TCI $1,112
Administrative
Charges 2% of TCI $2,224
Capital Recovery 15 yrs @ 10% $14,621
TOTAL INDIRECT ANNUAL COST (TIAC) $33,582
RECOVERY CREDIT (RC) $23,197
TOTAL ANNUAL COST TDAC + TIAC - RC $38,912
(TAG)
aJuly 1989 dollars
^Assumes refrigerated condenser system operated continuously
and required 0.5 hour of labor per 8-hour shift.
C-32
-------
APPENDIX D
APPLICATION OF STEAM STRIPPER COSTING METHODOLOGY
D.I INTRODUCTION
This appendix presents an application of the costing
methodology given in Section 3.2.3 of this volume to an
example steam stripper system. The example steam stripper
system is based on the configuration shown in Figure 2-8 of
this volume and is designed to treat an example wastewater
stream with a flow of 500 £pm (see Table D-l). Because the
cost of a steam stripper does not vary greatly with inlet
organic concentration, the wastewater composition is not
shown.
D.2 SYSTEM DESIGN
The first step in costing a steam stripper system is to
design the system.
D.2.1 Data Collection for Steam Stripper Design
Information on the design and operation of steam stripper
systems is available from studies conducted by EPA. During
the Industrial Wastewater Project, EPA obtained information on
approximately 15 steam strippers from facility responses to
CAA Section 114 information requests in 1987. Information
was also gathered during site visits to eight chemical
manufacturing facilities operating steam strippers to remove
organic compounds from wastewater. ••••'•• The EPA
also gathered data on steam stripper operation as a part of
the Hazardous Waste TSDF Project. During this project, data
were gathered through field testing at three steam
. . 11.12,13
strippers.
D-l
-------
TABLE D-l. DESIGN AND OPERATING BASIS FOR A 500-£pm
STEAM STRIPPING SYSTEM
1. Wastewater Flow: 500 £/min (132 gal/min)
2. Stripper Operating Period: 24 hr/day x 300 day/yr =
7,200 hr/yr
3. Wastewater Storage: Wastewater feed collection tank with
48-hr retention time
4. Steam Stripping Column:
Configuration: countercurrent flow, 9.0-m sieve
tray column (10 trays)
Steam Flow Rate: 0.096 kg of steam/i of waste feed
(0.8 Ib steam/gallon)
Wastewater Feed Temperature: 35 °C (95 °F)
Column Diameter: 0.98 m (3.2 ft)
Active Column Height: 6.5 m (21.3 ft)
Total Column Height: 9.0 m (29.5 ft)
Liquid Loading: 39,900 £/hr/m2 (979 gal/hr/ft2)
5. Condenser:
Configuration: Water-cooled
Primary Condenser Outlet Vapor
Temperature: 50 °C (122 °F)
6. Overhead Control: Vent to existing on-site combustion
or other control device
7. Bottoms Control: Feed to existing on-site Wastewater
treatment facility or POTW
D-2
-------
The EPA also gathered data on steam strippers during the
development of effluent guidelines for the OCPSF, Pesticide,
and Pharmaceutical Manufacturing industries. In response to
CWA Section 308 information requests, 63 OCPSF facilities
reported using a total of 108 steam strippers as in-plant
control for process wastewater.14 Field testing also
contributed data on steam strippers operating at three OCPSF
facilities.15' 6'17 Data available for steam strippers in use
at pesticide and pharmaceutical manufacturing industry
facilities came from eight pesticide facilities and eight
1S 19
pharmaceutical facilities. ' These strippers are similar
to strippers used in the SOCMI.
The steam stripper systems that are considered treat
wastewater streams that vary in flow rate and composition, and
some streams contain relatively high levels of suspended
solids. However, although the wastewater characteristics
vary, the basic steam stripper system shown in Figure 2-8 can
be designed and operated to achieve high efficiency in
removing organic compounds from most streams.
The organic removal performance of 13 steam stripper
systems was r asured during field tests sponsored by the EPA
and provided through CAA Section 114 questionnaire responses
and site visits to industrial facilities. Information
gathered during these efforts is supplied in Appendix C of the
Control Technology Center Document for Industrial Wastewater
Volatile Organic Compound Emissions—Background Information
for BACT/LAER Determinations. These data are summarized in
Table D-2. All significant figures have been retained for
intermediate values reported in this table to make it easier
for the reader to follow the calculation and to avoid
potential error due to round-off. It should not be inferred
that the intermediate values represent the actual number of
significant figures. The organic removals presented in
Table D-2 range from 76 percent for Site 7 to greater than
99.96 percent for Site C.
D-3
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D-6
-------
D.2.2 Development of Equipment Design and Operating
Parameters
The equipment design and operating parameters for the
example wastewater stream were developed through a design
evaluation performed using Advanced System for Process
Engineering (ASPEN) , a computer software program intended
for the rigorous design of distillation columns. The major
design parameters in the ASPEN steam stripper model are based
on field experience and published information. The diameter
was calculated assuming a velocity of 80-percent flooding
conditions. In addition, the following engineering
assumptions were made:
Operating pressure of one atmosphere;
Isothermal column operation;
Constant molal overflow (i.e., one mole of aqueous
phase vaporized for each mole of steam condensed);
and
• Linear equilibrium and operating equations (i.e.,
Henry's Law is valid for each organic compound at
the concentrations encountered in the stripping
column).
The design stripper contains 10 trays. A tray efficiency
of 80 percent was assumed to estimate the actual number of
stages for the column. A tray spacing of 0.50 m (1.64 ft) was
assumed to estimate the active column height. To approximate
the total column height, a total of 2.5 m (8.20 ft) of
nonactive entrance and exit column was assumed. The system
components are discussed in the following paragraphs.
The wastewater storage tank is sized to provide a desired
retention time of 48 hours for the stripper feed stream,
assuming 1500 m (4921 ft) of piping will be needed to combine
multiple batch and/or continuous streams for treatment by the
same steam stripper.
The remaining equipment in the steam stripper system was
designed using ASPEN. The steam stripper column is designed
as a sieve tray unit with countercurrent flow. The column is
operated at a typical steam-to-wastewater-feed ratio of
D-7
-------
0.096 kg of steam per t (0.80 Ib/gal) of wastewater. The
liquid loading of the column is 39,900 £/hr/m2
(979 gal/hr/ft2) .
Emphasis was placed on generating a design that would be
most cost effective, would be within practical design
parameters, and would remove virtually 100 percent of the
highly volatile compounds. (The controlling compound used for
design purposes was benzene.) A column height of 9 m (30 ft)
with a total of 10 sieve trays is used for the steam stripper
unit.
The overheads from the steam stripper are recovered with
a water-cooled condenser. The condenser is designed for an
outlet vapor temperature of 50 °C (122 °F) with an overall
heat-transfer coefficient of 1,000 j/m2/s/°K (0.05
Btu/ft /s/°F) . The organic phase of the overheads stream is
recovered from the overheads decanter. The overheads vapor
from the primary condenser is assumed to be vented to the feed
storage tank which is routed to an existing on-site combustion
device or other control device.
The bottoms from the steam stripper are fed to the
existing wastewater treatment facility. Before discharge from
the stripper system, the bottoms pass through a feed preheater
to enhance the efficiency of the steam stripper. The overall
heat-transfer coefficient used by ASPEN for the feed preheater
is 1,000 j/m2/s/°K (0.05 Btu/ft2/s/°F) .
Pumps are installed to transfer the wastewater from the
feed tank to the stripper, from the stripper to the
feed/bottoms heat exchanger, from the decanter to the
collection pot, and from the collection pot to storage.
There are five vents included in the system design for
the storage tank and decanter. The vapors from these vents
are ducted to a flare or other control device.
D.3 SYSTEM COSTING
Total capital investment and total annual cost for the
example steam stripper system are estimated using the
methodology described in Section 3.2.3 of this volume.
D-8
-------
D.3.1 Total Capital Investment
Steam stripper costs are estimated using the equipment
size generated by ASPEN. The purchase cost of each piece of
process equipment is determined from various costing
algorithms. Purchased equipment costs are then used to
estimate total installation costs, and finally total capital
investment.21 The cost-estimating techniques presented are
based upon the size or capacity of the equipment and are
derived from actual construction projects. Table D-3
summarizes the costing algorithms and Table D-4 summarizes the
estimated equipment costs calculated for each component, the
Table D-4 estimated size or capacity, and the reference or
information source used to obtain the cost estimate for
treating 500 £pm (132 gpm). The initial estimates were based
on the equipment costs for the year in which the textbook or
journal article was published. These costs were then adjusted
to July 1989 dollars using the Chemical Engineering fabricated
equipment index for the appropriate month and year. The cost
for each individual component was summed to yield the basic
equipment cost for the example wastewater stream.
The basic equipment cost was then used to calculate the
purchased equipment cost, direct installation costs, and
indirect installation costs according to the equations given
in Section 3.2.3 and summarized here in Table D-5. Table D-5
also shows the cost for each of these items.
D.3.2 Total Annual Cost
Total annual cost for the example steam stripper system
is based on the operating parameters developed using ASPEN,
and the methodology given in Section 3.2.3 of this volume.
Table D-6 summarizes the cost equations and the costs
calculated for utilities, labor, maintenance, and indirect
annual costs.
A credit was also calculated for the organics that are
recovered through steam stripping. Although there are several
options available for disposal or use of the recovered organic
stream, for this cost estimate it is assumed that the organics
D-9
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D-12
-------
TABLE D-5.
ESTIMATION OF TOTAL CAPITAL INVESTMENT FOR A
STEAM STRIPPING UNITa'b
Cost
Component
Direct Equipment Costs^
Base Equipment Cost (BEC)C
Pipingd
Instrumentaton
Sales Tax and Freight
Purchased Equipment Cost (PEC)e
Direct Installation Costs
Foundations and Supports
Electrical
Erection and Handling
Painting
Insulation
TOTAL DIRECT INSTALLATION COST
Indirect Installation Costs
Engineering and Supervision
Construction & Field Expense
Construction Fee
Startup and Testing
Contingency
TOTAL INDIRECT INSTALLATION COST
Cost
Factor
TABLE D-4
$36.83/m
($11.23/ft)
0.1*[BEC + Pipe]
0.08 * [BEC + Pipe]
12% of PEC
1% of PEC
40% of PEC
1% of PEC
1% of PEC
10% of PEC
10% of PEC
10% of PEC
1% of PEC
3% of PEC
Component
Cost"
$284,000
$58,000
$34,000
$30,000
$49,000
$4,000
$162,000
$4,000
$4,000
$41,000
$41,000
$41,000
$4,000
$12,000
Total Capital Cost
Investment Reference
28
29
29
$406,000
29
29
29
29
29
$223,000
29
29
29
29
29
$140,000
TOTAL CAPITAL INVESTMENT
$768,000
aJuly 1989 dollars.
"Based on 500 Ppm wastewater flow.
CBEC = Base equipment cost.
"Additional piping for combination of five wastewater streams is assumed to total approximately 1,500 m
(4,900 ft). Vapor vent lines required for storage tanks and decanters. Each vent line was assumed to be
11 m (36 ft) in length and constructed of 5.1-cm (2-in) diameter schedule 40 steel pipe.
ePEC = Purchased equipment cost.
D-13
-------
TABLE D-6. ESTIMATION OF TOTAL ANNUAL COST
FOR A STEAM STRIPPING UNITa'b
Cost
Component
Direct Annual Costs
Utilities
Electricity
Steam
Water
Labor
Operating Labor
Supervision & 15%
Administration
Maintenance
Labor
Materials 100%
TOTAL DIRECT ANNUAL COST
Indirect Annual Costs
Overhead
Property Taxes
Insurance
Administrative Charges
Capital Recovery
TOTAL INDIRECT ANNUAL COST (TIAC)
RECOVERY CREDIT
TOTAL ANNUAL COST
ANNUAL WASTE THROUGHPUT
COST PER UNIT WASTEWATER ($/M6)
COST PER LITER WASTEWATER FEED {$/!)
Cost Annual
Factor Consumption
$0.0509/kWhr 78,000 kWhrc
$7.68/Mg 29,000 Mgd
($7.77/ton) (32,000 tons)
$0.058/1,000 « 780,000,000 Je
($0.22/1,000 gal) (206,900,000 gal)
$13.20/hr 450 hrs
of Operating Labor
$14.50/hr 450 hrs
of Maintenance Labor
60% of All Labor
and Materials
1% of TCIf
1% of TCI
2% of TCI
10% @ 15 yrs
TDAC + TIAC - RCh
220,000 Mg/yr
TAC/AWT1
TAC/FLOWJ 260,000,000 i/yr
Annual
Cost
$4,000
$220,000
$45,000
$6,000
$1,000
$7,000
$7,000
$290,000
$13,000
$8,000
$8,000
$15,000
$101,000
$145,000
$17,0009
$418,000
$1.90
$0.0016
Cost
Reference
30
30
31
29
29
29
29
29
29
29
29
aJuly 1989 dollars.
"Based on 500 Ppm (132 gpm) wastewater flow.
C260 kWhr/day, 300 days/yr.
d97,000 kg/day (107 tons/day), 300 days/yr.
e2,600,000 {/day (690.000 gpd), 300 days/yr.
TCI = Total capital investment
^Recovery credit based on approximately 28,000 KJ/Kg (12,000 Btu/lb) heating value.
hTDAC + TIAC = RC = Total direct annual cost + Total indirect annual cost - Recovery credit.
1 TAG/AWT = Total annual cost per annual waste throughput.
JTAC/FLOW = Total annual cost per flow.
D-14
-------
will be used as fuel for an existing boiler. A heating value
of approximately 28,000 KJ/Kg (12,000 Btu/lb) was developed to
estimate the recovery credit in this example.
D-15
-------
D.4 References
1. U.S. Environmental Protection Agency, Office of Air
Quality Planning and Standards. Control Technology
Center Document for Industrial Wastewater Volatile
Organic Compound Emissions — Background Information for
BACT/LAER Determinations. EPA 450/3-90-004. Research
Triangle Park, NC. January 1990. pp. 4-4 to 4-5.
2. Letter and attachments from Radian Corporation to P.E.
Lassiter, P.E., EPA/OAQPS. July 17, 1987. The summary
of data provided in confidential Section 114
questionnaire responses.
3. Trip Report. Howie, R.H. and Vancil, M.A., Radian
Corporation, to file. May 12, 1987. Visit to Allied
Fibers.
4. Trip Report. Herndon, D.J. and Buchanan, S.K., Radian
Corporation, to file. May 22, 1987. Visit to Rhone-
Poulenc/Agricultural (RP Ag) Company.
5. Trip Report. Herndon, D.J. and Buchanan, S.K., Radian
Corporation, to file. September 2, 1987. Visit to
Fritzsche Dode & Olcott.
6. Trip Report. Herndon, D.J., Radian Corporation, to file.
May 6, 1987. Visit to PPG Industries.
7. Trip Report. Herndon, D.J. and Buchanan, S.K., Radian
Corporation, to file. May 21, 1987. Visit to Mobay
Chemical Company.
8. Trip Report. Herndon, D.J., Radian Corporation, to file.
May 4, 1987. Visit to Borden Chemical Company.
9. Trip Report. Herndon, D.J., Radian Corporation, to file.
May 5, 1987. Visit to Union Carbide Corporation.
10. Trip Report. Herndon, D.J., Radian Corporation, to file.
May 8, 1987. Visit to Dow Chemical Company.
11. U. S. Environmental Protection Agency. Hazardous Waste
Treatment, Storage, and Disposal Facilities (TSDF) -
Background Information for Proposed RCRA Air Emission
Standards, Volume 2: Appendices. Preliminary Draft.
Research Triangle Park, NC. March 1988. pp. F-146 to F-
149.
12. Ref. 11, pp. F-151 to F-147.
13. Ref. 11, pp. F-151 to F-155.
D-16
-------
14. Memorandum from Herndon, D. J., Radian Corporation, to
Industrial Wastewater file. May 20, 1988. Summary of
facilities reporting use of steam stripping.
15. Environmental Science & Engineering, Inc., and SAIC.
Plant No. 4 - Organic Chemicals Best Available Technology
Long-Term Field Sampling. Prepared for U. S.
Environmental Protection Agency, Office of Water
Regulations and Standards. July 1985.
16. Ref. 15, Plant No. 7.
17. Ref. 15, Plant No. 15.
18. Code of Federal Regulations. Title 40, Part 455.
Pharmaceutical Chemicals Category Effluent Limitations
Guidelines, Pretreatment Standards, and New Source
Performance Standards. Washington, DC. U.S. Government
Printing Office. October 4, 1985.
19. Code of Federal Regulations. Title 40, Part 439.
Pharmaceutical Manufacturing Point Source Category
Effluent Limitations Guidelines, Pretreatment Standards,
and New Source Performance Standards. Washington, DC.
U.S. Government Printing Office. October 27, 1983.
20. Advanced System for Process Engineering (ASPEN).
Massachusetts Institute of Technology for the Department
of Energy. DOE/MC/16481-(3 vols) 1201, 1202, 1203.
1981.
21. U. S. Environmental Protection Agency. Office of Air
Quality Planning and Standards. OAQPS Control Cost
Manual. Chapter 2: Cost Estimating Methodology.
4th Edition. EPA-450/3-90-006. Research Triangle Park,
N.C. January 1990. pp. 2-5 to 2-8.
22. Corripio, A.B., K.S. Chrien, and L.B. Evans. Estimate
Costs of Heat Exchangers and Storage Tanks via
Correlations. Chem.Eng. p. 145. January 25, 1982.
23. Peters, M.S., and K.D. Timmerhaus. Plant Design and
Economics for Chemical Engineers. Third Ed. New York,
McGraw-Hill Book Company. 1980. pp. 768-773.
24. Corripio, A.B., A. Mulet, and L.B. Evans. Estimate Costs
of Distillation and Absorption Towers via Correlations.
Chem.Eng. December 28, 1981. pp. 180-82.
25. Hall, R.S., W.M. Vatavuk, J. Matley. Estimating Process
Equipment Costs. Chem.Eng. November 21, 1988.
pp. 66-75.
26. Ref. 23, p. 572, Figure 13-58.
D-17
-------
27. Telecon. Gitelman, A., Research Triangle Institute with
Hoyt Corporation. Cost of flame arrestors. September 9,
1986.
28. Richardson Process Plant Construction Estimation
Standards: Mechanical and Electrical. Volume 3,
Richardson Engineering Services, Inc., Mesa, AZ, 1988.
pp. 15-40.
29. Vatavuk, W.M., and R.B. Neveril. Part II: Factors for
Estimating Capital and Operating Costs. Chem.Eng.
November 3, 1980. pp. 157-182.
30. Memorandum from Peterson, P., Research Triangle
Institute, to Thorneloe, S., EPA/OAQPS. January 18,
1988. Basis for steam stripping organic removal
efficiency and cost estimates used for the source
assessment model (SAM) analysis.
31. Ref. 2, p. 4-27.
32. Memorandum from Watkins, S., Radian Corporation, to
Lassiter, P., EPA/CPB. January 20, 1992. Development
of a recovery credit for organic compounds removed from
wastewater streams by the design steam stripper for the
HON.
D-18
-------
APPENDIX E
EXAMPLE COSTS FOR THE INSTALLATION OF
AN INTERNAL FLOATING ROOF IN
A FIXED ROOF TANK
This appendix presents a set of example calculations for
estimating the total annual cost for the installation and
operation of an internal floating roof in an example fixed
roof tank.
The purpose of this appendix is to demonstrate the
approach used in the HON analysis. In the calculations below,
all significant figures have been retained until the final
calculation to make it easier for the reader to follow the
calculation and to avoid potential error due to round off of
intermediate calculations. It should not be inferred that the
intermediate results represent the actual number of
significant figures.
The example tank stores styrene and has a capacity of
2,000,000 gallons. Additional design parameters for the
example tank are presented in Table E-l of this appendix. The
internal floating roof, installed in this example tank, will
have a liquid-mounted primary seal and controlled deck
fittings.
1. Before an internal floating roof can be installed in
the example tank, the tank must be cleaned and degassed. The
cost for cleaning and degassing a tank is based on its
capacity.
Cstdegas = 7.61 ($/gal) * [tank_size, gal]0-5132
Cstdegas = 7.61 * [2,000,000 gal]0-5132
Cstdegas = $13,034 per tank
E-l
-------
TABLE E-l. MODEL TANK DESIGN PARAMETERS
Parameter Description
Tank capacity (tank size)
Annual tank throughput
Tank orientation and type
Tank diameter (tank_dia)
Tank height
Average tank vapor space
height
Adjustment factor for small
diameter tanks
Tank paint factor - white
roof and aluminum color shell
Average ambient diurnal
temperature change
Product factor - organic
liquid other than crude oil
Stored product - HAP
Product molecular weight
Product specific gravity
Product vapor pressure at
25 °C
Product Antoine coefficients
A
B
C
Product average market price
Units
gal
gal/yr
feet
feet
feet
dimension less
dimension less
OF
dimensionless
Ib/lb mol
mmHg
psia
dimensionless
dimensionless
dimensionless
$/kg
Value
2,000,000
18,057,775
vertical
fixed roof
85
47
23.6
1.0
1.3
20
1.0
styrene
104.16
0.906
6.591
0.12748
7.140
1574.51
224.09
0.73
E-2
-------
2. Determine the installed capital cost for a new
internal floating roof having a primary liquid-mounted seal
and controlled deck fittings.
Cstroof = [509 ($/ft) * tank_dia(ft)] + 1160 ($)
Cstroof = [509 ($/ft) * 85 (ft)] + 1160 ($)
Cstroof = $44,425
3. The total capital investment (TCI) for the
installation of an internal floating roof in the example tank
is the sum of the tank preparation costs and the installed
capital cost for the roof.
Total , , . Installed
Tank cleaning
capital , capital
= and +
investment . cost for
degassing cost
(TCI) * * floating roof
TCI (July 1989$) =$13,034 + $44,425
TCI (July 1989$) = $57,459
4. Indirect annual costs (ann_fix) are estimated as a
percentage of the total capital investment (TCI) and include
capital recovery; maintenance charges (5 percent); inspection
charges (1 percent); and taxes, insurance, and administrative
charges (4 percent). The capital recovery factor (0.163) is
based on 10 percent interest over 10 years.
ann_fix (July 1989$) = $57,459 * [0.163 + 0.05 + 0.01 + 0.04]
Indirect Annual Cost (July 1989$) = $15,112/yr
5. The direct annual cost (ann__var) is a cost savings
equal to the value of the recovered product. This product
recovery credit is based on the HAP emission reduction
achieved by the internal floating roof and the market value of
the HAP.
E-3
-------
Direct Annual
Cost
(July 1989 $)
HAP emissions
reduction
(Mg/yr)
-Average market
value of HAP
($/Mg)
ann_var =5.17 (Mg/yr) * [-0.73 ($/kg)] * 1,000 (kg/Mg)
Direct Annual Cost (July 1989$) = -$3,774/yr
6. The total annual cost (TAG) for the installation and
operation of an internal floating roof is the sum of the
indirect annual cost and direct annual cost.
Total annual cost = Indirect annual cost + Direct annual cost
Total annual cost = $15,112 + [-$3,774]
Total annual cost = $ll,338/yr
E-4
-------
APPENDIX F
EXAMPLE COST METHODOLOGY FOR EQUIPMENT LEAKS
The purpose of these calculations is to demonstrate the
methodology used for calculating the cost of controlling
emissions from equipment leaks. Note that the sources of all
data used in the calculations are documented in the memorandum
"Final Cost Impacts Analysis for HON Equipment Leaks,"
April 15, 1991, contained in the equipment leaks docket. The
methodology is first described and then applied to an example
plant. Note that the costs presented reflect estimated costs
associated with implementing control as required in the
negotiated regulation. These costs may not be applicable to
other potential equipment leak standards nor to any individual
process unit. The following steps summarize how the costs are
calculated.
1. Determine a base cost for each type of component for
which a control equipment requirement is set by the
negotiated regulation. Cost are based on
information collected from vendors and published
sources. Convert all costs to July 1989 dollars.
2. Determine capital costs for process units by
multiplying the equipment count by the base cost for
each type of component. To this, add the purchase
cost of one monitoring instrument for each process
unit. (Note that process units controlled at
baseline will not incur capital costs for monitoring
instruments and caps for open-ended lines.)
3. Determine initial monitoring and leak repair cost
estimates for uncontrolled process units, based on
F-l
-------
equipment counts, initial leak frequencies, and
subsequent repair requirements.
4. Determine annual leak detection and repair cost
estimates for all process units based on equipment
counts, monitoring frequencies, leak frequencies,
and subsequent repair requirements.
5. Determine the annualized capital costs by
multiplying the capital cost values for each process
unit by the capital recovery factor which is
determined based on interest rate and equipment
life.
6. Determine annual costs for maintenance and
miscellaneous charges associated with equipment leak
control.
7. Determine the total annual cost by subtracting the
recovery credit from other annual costs (annualized
capital costs and annual operating costs). The
recovery credit is determined by multiplying the
annual emission reduction by the average VOC cost.
F.I BASE COSTS DEVELOPMENT
The following equations estimate base costs for control
equipment including a monitoring instrument, replacement pump
seals, compressors, pressure relief devices, open ended lines,
and sample connections. Chemical Engineering Plant Cost Index
values are used throughout the base cost development to
convert all costs to July 1989 dollars. A summary of the base
costs for control equipment is presented in Table F-l. Note
that total values for control equipment may not exactly add up
due to rounding errors.
Monitoring Instrument:
The price of a monitoring instrument is approximately $6,500
in July 1989 dollars.
Replacement Pump Seals:
The fourth quarter 1978 cost of a replacement seal for a pump,
including a 50 percent credit for the old seal, was $113.
F-2
-------
TABLE F-l. BASE COSTS FOR CONTROL EQUIPMENT
Base Cost
(per component)
Equipment Type (July 1989 $)
Monitoring Instrument 6,500
Replacement Pump Seal 180
Compressor 6,240
Pressure Relief Device:3
Rupture Disks 78
Holders, Valves, Installation, etc. 3852
Open-Ended Lines 102
Sample Connections 408
aPressure Relief Device costs are split into two portions
because Rupture Disks have a 2-year equipment life and all
other pieces (Holders, Valves, etc.) have a 10-year life.
F-3
-------
Using the averaged Chemical Engineering Plant Cost Index Ratio
of (356.0/224.7] to convert to July 1989 dollars:
$113 * (356.0/224.7) = $180.
Compressors (Closed-Vent System);
A closed vent system for a compressor consists of piping, plug
valves, and a flame arrestor. The costs for equipment and
installation of a closed-vent system are described in the
following steps.
1. 122 m of 5.1 cm diameter schedule 40 steel pipe.
Piping costs are $261.03/100 ft in 1988 dollars.
Using the Chemical Engineering Plant Cost Index
Ratio of (356.0/342.5) to convert to July 1989
dollars:
(122 m) * (3.28 ft/m) * ($261.03/100 ft) *
(356.0/342.5) = $1090.
2. Installation of this piping will require 66 hrs of
labor at a rate of $22.50/hr:
(66 hrs) * ($22.50/hr) = $1490.
3. Three 5.1 cm cast steel plug valves are required at
a 1988 cost of $783.75 each. Use the Chemical
Engineering Plant Cost Index Ratio of (356.0/342.5)
to convert to July 1989 dollars.
3 * ($783.75) * (356.0/342.5) = $2,440
4. Installation of these valves will require 2 hours of
labor at a rate of $22.50/hr:
3 * ($22'. 50/hr) * (2 hrs) = $135.
5. One metal gauze flame arrestor will cost $869 based
a 1990 vendor quote. Use the Chemical Engineering
Plant Cost Index Ratio of (356.0/358.7) to convert
to July 1989 dollars:
($869) * (356.0/358.7) = $860.
6. Installation of the flame arrestor will require 10
hrs of labor at a rate of $22.50/hr:
($22.50/hr) * (10 hrs) = $225.
Total Compressor Closed Vent Cost:
$1,090 + $1,490 + $2,440 + $135 + $860 + $225 = $6,240.
F-4
-------
Pressure Relief Devices (Rupture Disk Assembly):
A rupture disk assembly for a pressure relief device consists
of a rupture disk, rupture disk holder, pressure gauge, bleed
valve, gate valve, offset, and relief valve. The costs for
equipment and installation of a rupture disk assembly for a
pressure relief device are as follows:
1. Three 7.6 cm stainless steel rupture disks cost $235
in 1990. Use the Chemical Engineering Plant Cost
Index Ratio of (356.0/358.7) to convert to July 1989
dollars:
[($235)/(3)] * (356.0/358.7) = $78.
2. One 7.6 cm carbon steel rupture disk holder cost
$583 in 1990. Use the Chemical Engineering Plant
Cost Index Ratio of (356.0/358.7) to convert to
July 1989 dollars:
($583) * (356.0/358.7) = $580.
3. One 0.6 cm pressure gauge with dial face cost $15 in
last quarter 1978. Use the Chemical Engineering
Plant cost Index Ratio of (356.0/224.7) to convert
July 1989 dollars:
($15) * (356.0/224.7) = $24.
4. One 0.6 cm carbon steel bleed valve cost $25 in last
quarter 1978. Use the Chemical Engineering Plant
Cost Index Ratio of (356.0/224.7) to convert to
July 1989 dollars:
($25) * (356.0/224.7) = $40.
5. Installation of the pressure gauge and bleed valve
will require 16 hours of labor at a rate of
$22.50/hr:
(16 hours) * ($22.50/hr) = $360.
6. One 7.6 cm gate valve cost $385.60 in 1988. Use the
Chemical Engineering Plant Cost Index Ratio of
(356.0/342.5) to convert to July 1989 dollars:
($385.60) * (356.0/342.5) = $400
7. Installation of the gate valve will require
1.5 hours of labor at a rate of $22.50/hr:
(1.5 hours) * ($22.50/hr) = $34
F-5
-------
8. An offset consisting of one 10.2 m tee and one
10.2 cm elbow cost $47.36 in 1988. Use the Chemical
Engineering Plant Cost Index Ratio of (356.0/342.5)
to convert to July 1989 dollars:
($47.36) * (356.0/342.5) = $49.
9. Installation of the offset will require 8.5 hours of
labor at a rate of $22.50/hr:
(8.5 hours) * ($22.50/hr) = $191.
10. Retrofitting of existing pressure relief valves will
require a new 7.6 cm stainless steel body and trim
relief valve which cost $1230.80 in the last quarter
of 1978. Use the Chemical Engineering Plant Cost
Index Ratio of (356.0/224.7) to convert July 1989
dollars:
($1230.80) * (356.0/224.7) = $1,950.
11. Installation of the relief valve will require
10 hours of labor at a rate of $22.50/hr:
($22.50/hr) * (10 hrs) = $225.
Total Pressure Relief Device Cost:
$78 + $580 + $24 + $40 + $360 + $400 + $34 + $49 + $191 +
$1950 + $225 = $3,930.
Open-Ended Lines;
Control equipment for an open ended line consists of one
2.5 cm gate valve which cost $67.50 in 1988. Use the Chemical
Engineering Plant Cost Index Ratio of (356.0/342.5) to convert
to July 1989 dollars:
($67.50) * (356.0/342.5) = $70.
Installation of the gate valve will require 1.4 hours of labor
at a rate of $22.50/hr:
(1.4 hours) * ($22.50/hr) = $32.
Total Open-Ended Line Cost = $70 + $32= $102.
Sample Connection:
A closed purge sample connection system consists of piping and
three ball valves. The costs for equipment and installation
are as follows:
F-6
-------
1. 6 m of 2.5 cm diameter schedule 40 steel pipe cost
$120.64/100 ft in 1988. Use the Chemical
Engineering Plant Cost Index Ratio of (356.0/342.5)
to convert to July 1989 dollars:
(6 m) * (3.28 ft/m) * ($120.64/100 ft) *
(356.0/342.5) = $25.
2. Installation of the piping will require 1 hour of
labor at a rate of $22.50/hr:
(1 hour) * ($22.50/hr) = $23.
3. One 2.5 cm carbon steel ball valve cost $85.60 in
1988. Use the Chemical Engineering Plant Cost Index
Ratio of (356.0/342.5) to convert to July 1989
dollars:
(3) * ($85.60) * (356.0/342.5) = $265.
4. Installation of the three valves will require
1.4 hrs of labor each at a rate of $22.50/hr:
(3) * (1.4 hours) * ($22.50/hr) = $95.
Total Sample Connection Cost:
$25 + $23 + $265 + $95 = $408
F.2 CAPITAL COST DEVELOPMENT
Two sets of equations are used to develop capital costs
for the process units, one for controlled units and one for
uncontrolled units. Note that open-ended lines are routinely
controlled in all process units. If an open-ended line is
already controlled with a second valve, blind flange, or cap,
then capital costs for that open-ended line are zero dollars.
These equations are presented below.
Controlled Units
(Compressor Count) * ($6,240) = Compressor
Capital
(Pressure Relief Device Count * ($78) = Rupture Disk
Capital
(Pressure Relief Device Count) * ($3,852) = Holder,
Valve, Etc.
Capital
F-7
-------
(Sample Connection Count) * ($408)
Sum of above
Uncontrolled Units
(1 Monitoring Instrument) * ($6,500)
(Compressor Count) * ($6,240)
(Pressure Relief Device Count) * ($78)
(Pressure Relief Device Count) * ($3,852)
(Open-Ended Line Count) * ($102)
(Sample Connection Count) * ($408)
Sum of above
Sample
Connection
Capital
Total
Capital for
Controlled
Unit
Monitoring
Instrument
Capital
Compressor
Capital
Rupture Disk
Capital
Holder,
Valve, Etc.
Capital
Open-Ended
Line Capital
Sample
Connection
Capital
Total
Capital for
Uncontrolled
Unit
F.3 INITIAL LEAK DETECTION AND REPAIR COSTS
Leak detection and repair costs are estimated for
uncontrolled and controlled process units. Uncontrolled
process units have cost associated with implementing the leak
detection and repair program. Subsequent to implementing the
leak detection and repair program annual costs for
uncontrolled and controlled process units are equal.
For process units which are uncontrolled, the costs for
initial monitoring and leak repair valves in gas/vapor
service, valves in light liquid service, pumps in light liquid
F-8
-------
service, and connectors are estimated by the following
procedure.
Monitoring Costs;
Monitoring costs are based on a subcontractor's monitoring fee
of $2.50/component/monitoring event. This cost includes a
first attempt at repair.
(Equipment Type Count) * ($2.50/component) = Monitoring Cost
This is reported for each type of equipment covered.
Initial Repair Cost:
Determination of the cost of initial repairs requires three
steps. These steps are to determine the total number of leaks
found, determine the number of leaks requiring repair, and
determine the cost of repairs. Each step is described below.
1. The total number of leaks found during monitoring is
estimated by multiplying the number of components of
a specific type by the initial leak frequency. The
initial leak frequencies are:
Valves in gas/vapor service = 0.114
Valves in light liquid service = 0.065
Pump seals in light liquid service =0.20
Connectors = 0.021
(Equipment Type Count) * (initial leak frequency) =
Total number of leaks
2. The number of leaks requiring further repair (beyond
efforts made by the monitoring team) is determined
by multiplying the number of leaks found by the
percent of leaks requiring further repair. For
valves and connectors, this is 25 percent. For
pumps this is 75 percent. For example:
(Pump leaks) * (0.75) = Number of pumps requiring
further repair
F-9
-------
3. The cost of further repairs is determined for each
equipment type by the time required for each repair
times the labor rate of $22.50/hr. The time
required for each repair is:
Valves = 4 hours/repair
Pumps = 16 hours/repair
Connectors = 2 hours/repair
(pump^e^iring) * (l6 hrs ) * ($22-50) » Pun* repair
(further repair J \ repair/ \ hr J cost
The total cost for initial monitoring and repair is the sum of
monitoring costs and repair costs.
Total Cost for
(S Monitoring Costs) + (S Repair Costs) = initial monitoring
and repair
F.4 ANNUAL LEAK DETECTION AND REPAIR COSTS
For all process units, the costs for annual leak
detection and repair (LDAR) are estimated by the following
procedure.
Monitoring Costs;
Monitoring costs are based on a subcontractor's monitoring fee
of $2.00/component/monitoring event for every monitoring after
the first. This cost includes a first attempt at repair by
the monitoring team. The price is lower than the initial
monitoring fee because of reduced time requirements and lower
leak frequency.
(Equipment Type Count) * (Monitoring Frequency) *
($2.00/component/event) = Instrument Monitoring Cost
For pumps, an additional monitoring fee is added to cover
weekly visual inspections. This cost is calculated by
multiplying the number of pumps by the inspection rate of 1
inspection every 0.5 minutes times the frequency of 52 visual
monitoring/year, and then times the labor rate of $22.50/hr.
(Pump count) * (0.5 min/inspection) * (1 hr/60 min) *
(52 monitoring/yr) * ($22.50/hr) = Visual Monitoring Cost
F-10
-------
Repair Costs;
The annual leak repair costs are determined using three steps:
determine the total number of leaks found, determine the total
number of leaks requiring repair, and determine the cost of
repairs. These steps are discussed below.
1. The total number of leaks found during monitoring is
estimated by multiplying the number of components of
a specific type by the number of times the type of
equipment is monitored annually times the initial
leak frequency. The leak frequencies were taken
from the negotiated regulation as the maximum value
which did not require the implementation of a
quality improvement plan. These leak frequencies
are:
Valves in gas/vapor service =0.02
Valves in light liquid service = 0.02
Pump seals in light liquid service =0.10
Connectors = 0.005
The monitoring frequencies are as follows:
Valves in gas/vapor service = monthly3
Valves in light liquid service = monthly3
Pumps in light liquid service = monthly
Connectors = annually
(Equipment Count) * (Monitoring Frequency) *
(Leak Frequency) = Total number of leaks
2. The number of leaks requiring further repair (beyond
efforts made by the monitoring team) is determined
by multiplying the total number of leaks found by
the percentage of leaks requiring further repair.
For valves and connectors this is 25 percent. For
pumps this was determined to be 75 percent.
3. The cost of further repairs is estimated using the
same approach as described for the initial repair
costs.
aValves are monitored monthly in this example, but if less
than 2 percent are leaking, can be monitored quarterly.
F-ll
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The total cost for annual leak detection and repair is the sum
of monitoring and repair costs, plus administrative and
support cost. Administrative and support cost are estimated
to be 40 percent of the monitoring and repair labor expenses:
Administrative
(Monitoring labor + Repair labor) * (0.40) = and Support
Costs
Note that the cost for monitoring pressure relief devices once
annually is included in the LDAR costs. Further repair of
this equipment is not included, however, because it is assumed
to be covered by plant maintenance costs not the LDAR program.
F.5 ANNUALIZED CAPITAL COSTS
The annual ized capital costs for the process units is
determined by multiplying the capital cost for each type of
equipment by the Capital Recovery Factor for that type of
equipment. The Capital Recovery Factors were calculated using
the following equation:
- i * (1 + i)n
-
where: i = interest rate, expressed as a decimal
n = economic life of the component, in years.
For all components, the interest rate is 10 percent or 0.10.
For pump seals and rupture disks, the economic life is 2
years. For monitoring instruments it is 6 years, and for all
other equipment it is 10 years. Then the CRF values are:
CRF: 2 years = 0.58
CRF: 6 years = 0.23
CRF: 10 years = 0.163
/Component Type\ /r,PT?v _ / Annual ized \
\ Capital Cost ) (^Kt!> ~ \Capital Costs/
For uncontrolled units, the costs for initial LDAR are
annualized. This is done by multiplying the total cost for
F-12
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labor for the initial leak detection and repair by 1.4 to
40 percent for administration and support, and then
multiplying by the CRF for 10 years of 0.163.
('Initial^
Leak
Repair
Cost
* (1.4) * (0.163) =
Annualized
initial LDAR
Labor, Admin.,
(and Support CostJ
The cost of replacement pump seals associated with the initial
LDAR are also annualized. It was calculated by multiplying
the number of pumps requiring further repair by the CRF for
2 years times the cost of $180/pump seal.
/'Number of
Pumps
Requring
Repair
* (0.58) *
$180
pump seal
/Annualized costl
= of pump seal
V replacement >
F.6 ANNUAL MAINTENANCE COSTS AND MISCELLANEOUS COSTS
In addition to annual leak detection and repair program
costs and annualized capital costs, there are annual
maintenance charges and annual miscellaneous fees.
Annual Maintenance Charges:
Annual maintenance charges include maintenance costs
associated with the monitoring instrument, control equipment
installed on compressors, pressure relief devices, open-ended
lines, and sampling connections, and replacement pump seals.
These costs are described below.
1. In the last quarter of 1978, the cost for all
materials and labor required for maintenance and
calibration of the monitoring instrument was $2,700.
Updating this cost to July 1989 dollars is done
using the Chemical Engineering Plant Cost Index
Ratio Of (356.0/224.7)
$2700 * (356.0/224.7) = $4,280
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2. For compressors, pressure relief devices, open-ended
lines, and sampling connections, the annual
maintenance charge was calculated as 5 percent of
the capital charge for the same equipment type.
(Capital Cost) * (0.05) = Annual Maintenance Charge
For this analysis, it was assumed that open-ended
lines were controlled at baseline and would
therefore incur no capital cost. However, in order
to estimate the maintenance and miscellaneous
charges associated with open-ended lines, it was
necessary to calculate a capital cost value for use
in the above equation.
3. Replacement of pump seals is considered an annual
maintenance cost. The cost is calculated by
multiplying the number of pumps found leaking as a
result of the annual LDAR program by $180/pump seal.
Annual Miscellaneous Costs:
The annual miscellaneous costs include taxes, insurance, and
administration. For the monitoring instrument, compressors,
pressure relief devices, open-ended lines, sampling
connections, and pump seals. The annual miscellaneous charges
are calculated as 4 percent of the capital cost.
(Capital Cost) * (0.04) = Annual Miscellaneous Charges
For replacement pump seals, the annual miscellaneous
charge was calculated as 80 percent of the annual maintenance
cost for pump seals.
F.7 TOTAL ANNUAL COSTS
Total annual costs include all annual cost described
previously plus recovery credit associated with emission
reductions. The recovery credit for VOC's which are not lost
to the air through equipment leaks is calculated by
multiplying the annual emission reduction by the average VOC
cost.
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/Annual Emission\ /Average VOC\ _ Recovery
\ Reduction / \Cost($/Mg)j Credit
Total annual cost for a process unit is determined by
subtracting the Recovery Credit from other annual costs.
Other annual costs include the sum of the annualized capital
cost, and annual operating expenses for the LDAR program,
maintenance charges, and miscellaneous costs.
(AloLlifi1t-:LaZied) x /Operating\ /Recovery\ _ Total Annual
[ Costs J IE*Penses) " \ Credit ) ~ Cost
F.8 EXAMPLE CALCULATIONS
Example calculations are provided for estimation of
annualized capital costs and operating costs for uncontrolled
and controlled hypothetical process units. Total annual costs
are not calculated in these examples. BID Volume 1C,
Appendix E demonstrates calculation of total annual costs.
Note that all summations are rounded to 100's of dollars.
The following calculations are for an uncontrolled
process unit which has the following hypothetical equipment
counts:
Valves in gas/vapor service 414
Valves in light liquid service 1,179
Pumps in light liquid service 40
Connectors 2,662
Pressure Relief Devices (PRVs) 45
Compressors 2
Open-Ended Lines (OELs) 141
Sample Connection 35
Capital Costs
Monitoring Instrument = $ 6,500
Compressors
(2 compressors) * ($6,240/compressor) = $ 12,480
Rupture Disk Assemblies:
- Disks
(45 PRVs) * ($78/PRV) = $ 3,510
F-15
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- Holders, valves, etc.
(45 PRVs) * ($3,852/PRV)
Sample Connections
(35 sample con.) * ($408/sample con.)
Total Capital Cost
$173,340
$ 14.280
$210,100
Initial Monitoring and Leak Repair
Monitoring
Gas/Vapor Valves:
(414 valves) * (2.5O/component)
Light Liquid Valves:
(1,179 valves) * (2.5O/component)
Light Liquid Pumps:
(40 pumps) * ($2.50/component)
Connectors:
(2,662 connectors) ($2.50/component)
Initial Repair, Leaks Found:
Gas/Vapor Valves:
(414 valves) * (0.114 leaks/valve)
Light Liquid Valves:
(1,179 valves) * (0.065 leaks/valve)
Light Liquid Pumps:
(40 pumps) * (0.20 leaks/pump)
Connectors:
(2,662 connectors) *
(0.021 leaks/connectors)
Leaks Requiring Repair:
Gas/Vapor Valves: (47.2 leaks) * (0.25)
Light Liquid Valves: (76.6 leaks) * (0.25)
Light Liquid Pumps: (8.0 leaks) * (0.75)
Connectors: (55.9 leaks) * (0.25)
Repair Costs:
Gas/Vapor Valves:
(11.8 leaks) * (4 hrs/leak) * ($22.5/hr)
Light Liquid Valves:
(19.2 leaks) * (4 hrs/leak) * ($22.5/hr)
Light Liquid Pumps:
(6.0 leaks) * (16 hrs/leak) * ($22.5/hr)
Connectors:
(14.0 leaks) * (2 hrs/leak) * ($22.5/hr)
$ 1,035
$ 2,948
$ 100
$ 6,655
= 47.2 leaks
= 76.6 leaks
= 8.0 leaks
= 55.9 leaks
11.8 leaks
19.2 leaks
6.0 leaks
14.0 leaks
Total Initial LDAR Costs
Gas/Vapor Valves:
Light Liquid Valves:
$1,035 + $1,062
$2,948 + $1,728
$ 1,062
$ 1,728
$ 2,160
$ 630
$ 2,097
$ 4,676
F-16
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Light Liquid Pumps:
Connectors:
$ 100 + $2,160
$6,655 + $ 630
(12 times/yr) *
Annual Leak Detection and Repair
Monitoring Costs:
Gas/Vapor Valves:
(414 valves) * (12 times/yr) *
($2/time/valve)
Light Liquid Valves:
(1,179 valves) *
($2/time/valve)
Light Liquid Pumps:
(40 valves) * (12 times/yr) *
($2/time/valve)
Connectors:
(2,662 valves) * (1 time/yr) *
($2/time/concentration)
Pressure Relief Leak:
(45 PRVs) * (l time/yr) * ($2/time/PRV)
Pumps (Visual):
(40 pumps) * (52 times/yr) *(0.5 min/time)
(1 hr/60 min) * ($22.5/hr)
Annual Leaks Found:
Gas/Vapor Valves:
(414 valves) * (12 times/yr) *
(0.02 leaks/valve)
Light Liquid Valves:
(1,179 valves) * (12 times/yr) *
(0.02 leaks/valve)
Light Liquid Pumps:
(40 pumps) * (12 times/yr) *
(0.10 leaks/pump)
Connectors:
(2,662 connectors) * (1 time/yr) *
(0.005 leaks/connectors)
Leaks Requiring Repair:
Gas/Vapor Valves: (99.4 leaks) * (0.25)
Light Liquid Valves: (283.0 leaks) * (0.25)
Light Liquid Pumps: (48.0 leaks) * (0.75)
Connectors: (13.3 leaks) * (0.25)
$ 2,260
$ 7.285
$16,300
$ 9,936
$28,296
$ 960
$ 5,324
$ 90
$ 390
$45,000
99.4 leaks
283.0 leaks
48.0 leaks
13.3 leaks
24.9 leaks
70.8 leaks
36.0 leaks
3.3 leaks
F-17
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Cost of Repairs:
Gas/Vapor Valves:
(24.9 leaks) * (4 hrs/leak) * ($22.5/hr) = $ 2,241
Light Liquid Valves:
(70.8 leaks) * (4 hrs/leak) * ($22.5/hr) = $ 6,372
Light Liquid Pumps:
(36.0 leaks) * (16 hrs/leak) * ($22.5/hr) = $12,960
Connectors:
(3.3 leaks) * (2 hrs/leak) * ($22.5/hr) = $ 149
$21,700
Annualized Capital Charges
Monitoring Instrument: (0.23) * ($6,500) = $ 1,495
Compressors: (0.163) * ($12,480) = $ 2,034
Rupture Disk Assemblies:
-Disks: (0.58) * ($3,510) = $ 2,036
-Holders, Valves, etc: (0.163) * ($173,340) = $28,254
Sample Connections: (0.163) * ($14,280) = $ 2,328
Initial LDAR: (0.163) * ($16,300) * (1.4) = $ 3,720
Pump Seals: (6.0) * (0.58) * ($180) = $ 626
$40,500
Annual Maintenance Costs
Monitoring Instrument: = $ 4,280
Compressors: ($12,480) * (0.05) = $ 624
Rupture Disk Assemblies:
-Disks: ($3,510) * (0.05) = $ 176
-Holders, Valves, etc: ($173,340) * (0.05) = $ 8,667
Open-Ended Lines:a ($14,382) * (0.05) = $ 719
Sample Connections: ($14,280) * (0.05) = $ 714
Replacement Pump Seals:
(36.0 leaks) * ($180/seal) = $ 6.480
$21,700
Miscellaneous Annual Charges
Monitoring Instrument: ($6,500) * (0.04) = $ 260
Compressors: ($12,480) * (0.04) = $ 499
alf open-ended lines were uncontrolled, the capital cost to
control them would be 141 times $102 which equals $14,280.
As described in Section F.6, it is assumed that open-ended
lines are controlled at baseline, and the value for capital
cost for open-ended lines is calculated solely for the
purpose of estimating annual maintenance and miscellaneous
charges. This value for capital cost is not included in the
total capital cost for the process unit ($210,100).
F-18
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Rupture Disk Assemblies:
-Disks: ($3,510) * (0.04) = $ 140
-Holders, Valves, etc: ($173,340) * (0.04) = $ 6,934
Pressure Relief Devices =
Open-Ended Lines:a ($14,382) * (0.04) = $ 575
Sample Connections: ($14,280) * (0.04) = $ 571
Replacement Pump Seals: (6,480) * (0.8) = $ 5.184
$ 14,200
Labor Charges
Monitoring Cost: = $ 45,000
Repair Costs: = $ 21,700
Administration and Support:
(0.4) * ($45,000 + $21,700) = $ 26.700
$ 93,400
Total Annual Costs
Annualized Capital Charges = $ 40,500
Annual Operating Cost: ($21,700 + $14,200) = $ 35,900
Annual Labor Costs = $ 93.400
$169,800
The following calculations are for a baseline controlled
process unit having identical equipment counts to those used
in the previous example calculation.
Capital Costs
Compressors:
(2 compressors) * ($6,240/compressor) = $ 12,480
Rupture Disk Assemblies:
-Disks: (45 PRVs) * (78/PRV) = $ 3,510
-Holders,Valves, etc: (45 PRVs) *
($3,852 PRV) = $173,340
alf open-ended lines were uncontrolled, the capital cost to
control them would be 141 times $102 which equals $14,280.
As described in Section F.6, it is assumed that open-ended
lines are controlled at baseline, and the value for capital
cost for open-ended lines is calculated solely for the
purpose of estimating annual maintenance and miscellaneous
charges. This value for capital cost is not included in the
total capital cost for the process unit ($210,100).
F-19
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Sample Connections:
(35 sample connectors) *
($408/sample connector) = $ 14.280
$203,600
Annual Leak Detection and Repair
The annual leak detection and repair costs for a baseline
controlled unit are exactly the same as for a baseline
uncontrolled unit. Total monitoring cost is $45,000, repair
cost $21,700, and administration and support cost $26,700.
Annualized Capital Charges
Compressors: (0.163) * ($12,480) = $ 2,034
Rupture Disk Assemblies:
-Disks: (0.58) * ($3,510) = $ 2,036
-Holders, Valves, etc: (0.163) * ($173,340) = $28,254
Sample Connections: (0.163) * ($14,280) = $ 2.328
$34,700
Annual maintenance and miscellaneous costs for a baseline
controlled process unit are the same as for a baseline
uncontrolled unit. Annual maintenance costs are $21,700 and
miscellaneous costs are $14,200.
Labor Charge Summary:
Monitoring Cost: = $ 45,000
Repair Costs: = $ 21,700
Administration and Support: . = $ 26.700
$ 93,400
Total Annual Costs
Annualized Capital Charges = $ 34,700
Annual Operating Cost: = $ 35,900
Annual Labor Costs = $ 93.400
$164,000
F-20
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