Control of Volatile Organic Compound Emissions from Manufacture of High-Density Polyethylene, Polypropylene, and Polystyrene Resins


                              EPA-450/3-83-008
          Guideline Series

   Control of Volatile Organic
   Compound Emissions from
  Manufacture of High-Density
Polyethylene,  Polypropylene, and
        Polystyrene  Resins
         Emission Standards and Engineering Division
         U.S. ENVIRONMENTAtpROTECTION AGENCY
            Office of Air, Nof^e, and Radiation
          Office of Air Quality Pfenning and Standards
         Research Triangle Park;.North Carolina 27711

                November 1983

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                                  GUIDELINE SERIES

The guideline series of reports is issued by the Office of Air Quality Planning and Standards
(OAQPS) to provide information to state and local air pollution control agencies; for example, to
provide guidance on the acquisition and processing of air quality data and on the planning and
analysis requisite for the maintenance of air quality. Reports published in this series will be
available - as supplies permit - from the Library Services Office (MD-35), U.S. Environmental
Protection Agency, Research Triangle Park, North Carolina 27711, orfor a nominal fee, from the
National Technical Information Service, 5285 Port Royal Road, Springfield, Virginia 22161.

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                           TABLE OF CONTENTS
1.0  INTRODUCTION
2.0  PROCESS AND POLLUTANT EMISSIONS  ....
     2.1  INTRODUCTION  	 ....
     2.2  POLYPROPYLENE
          2.2.1  General Industry Description
          2.2.2  Model Plant
     2.3  HIGH-DENSITY POLYETHYLENE
          2.3.1  General Industry Description
          2.3.2  Model Plant	
     2.4  POLYSTYRENE	
          2.4.1  General Industry Description
          2.4.2  Model Plant  	 ..
     2.5  REFERENCES FOR CHAPTER 2	
3.0  EMISSION CONTROL TECHNIQUES
     3.1  CONTROL BY COMBUSTION TECHNIQUES.  .
          3.1.1  Flares ...........
          3.1.2  Thermal Incinerators ....
          3.1.3  Catalytic Incinerators ...
          3.1.4  Industrial Boilers
     3.2  CONTROL BY RECOVERY TECHNIQUES
          3.2.1  Condensers	
          3.2.2  Adsorbers
          3.2.3  Absorbers
     3.3  REFERENCES FOR CHAPTER 3.  .....
4.0  ENVIRONMENTAL ANALYSIS OF RACT
     4.1  INTRODUCTION. .....
     4.2  AIR POLLUTION
     4.3  WATER POLLUTION	
     4.4  SOLID WASTE DISPOSAL. .......
     4.5  ENERGY	
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5.0  CONTROL COST ANALYSIS OF RACT	
     5.1  BASES OF COST ANALYSES	
          5.1.1  Thermal Incinerator Design
                   and Cost Basis 	
          5.1.2  Flare Design and Cost Basis	
          5.1.3  Catalytic Incinerator Design
                   and Cost Basis 	
          5.1.4  Condenser Design and Cost Basis. . .  .
     5.2  EMISSION CONTROL COSTS	
          5.2.1  Polypropylene (PP)	
          5.2.2  High-Density Polyethylene (HOPE) . .  .
          5.2.3  Polystyrene (PS) 	
     5.3  COST EFFECTIVENESS OF RACT	
     5.4  REFERENCES  FOR CHAPTER 5	
APPENDIX A - LIST OF  COMMENTERS 	
APPENDIX B - COMMENTS ON MAY 1982 DRAFT CT6 DOCUMENT.  .
APPENDIX C - MAJOR  ISSUES AND RESPONSES 	
     C.I  THE  INCLUSION OF FLARES AS RACT  	
     C.2  ACCEPTABILITY OF CONDENSERS, CATALYTIC
             INCINERATORS, ABSORBERS, AND
             PROCESS HEATERS AS RACT 	
     C.3  STRINGENCY  OF RACT	  .  .  .
     C.4  BASIS  OF  COST ANALYSIS	
          C.4.1   Origin of Costs  (GARD vs
                    Enviroscience	
          C.4.2   Cost Effectiveness Calculations.  .  .  .
          C.4.3   Miscellaneous	  .
     C.5  SCOPE  OF  CTG:  POLYSTYRENE CONTINUOUS
          PROCESS 	
     C.6  SCOPE  OF  THE CTG:   HIGH DENSITY POLYETHYLENE,
          LIQUID PHASE SOLUTION PROCESS AND OTHER
          PROCESSES NOT CURRENTLY INCLUDED	
 APPENDIX  D  - EMISSION SOURCE TEST DATA	
     D.I  FLARE  VOC EMISSION TEST DATA	,.
          D.I.I   Control  Device	  .   . .
           D.I.2   Sampling and Analytical  Techniques  .  .
           D.I.3   Test Results 	
5-1
5-1

5-2
5-11

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D-3
 D-3
D-5

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     D.2  THERMAL INCINERATOR VOC EMISSION TEST DATA. .  .
          D.2.1  Environmental  Protection Agency (EPA)
                   Polymers Test Data 	
          D.2.2  Environmental  Protection Agency (EPA)
                   Air Oxidation Unit Test Data 	
          D.2.3  Chemical  Company Air Oxidation
                   Unit Test Data 	
          D.2.4  Union Carbide Lab-Scale Test Data. . .  .
     D.3  VAPOR RECOVERY SYSTEM VOC EMISSION TEST DATA.  .
     D.4  DISCUSSION OF TEST RESULTS AND THE TECHNICAL
            BASIS OF THE POLYMERS AND RESINS VOC
            EMISSIONS REDUCTION REQUIREMENT .. 	
          D.4.1  Discussion of Flare Emission Test
                   Results	
          D.4.2  Discussion of Thermal Incineration
                   Test Results . .	
     D.5  REFERENCES FOR APPENDIX D 	
APPENDIX E - DETAILED DESIGN AND COST ESTIMATION
             PROCEDURES 	
     E.I  GENERAL 	
     E.2  FLARE DESIGN AND COST ESTIMATION PROCEDURE. .  ,
          E.2.1  Flare Design Procedure	
          E.2.2  Flare Cost Estimation Procedure	
     E.3  THERMAL INCINERATOR DESIGN AND COST ESTIMATION
          PROCEDURE 	  	
          E.3.1  Thermal Incinerator Design Procedure .  ,
          E.3.2  Thermal Incinerator Cost  Estimation
                   Procedure.	 .  .
     E.4  CATALYTIC INCINERATOR DESIGN AND COST
          ESTIMATION PROCEDURE	,
          E.4.1  Catalytic Incinerator Design Procedure .
          E.4.2  Catalytic Incinerator Cost Estimation
                   Procedure. ......  	 ,
     E.5  SURFACE CONDENSER DESIGN AND COST ESTIMATION
          PROCEDURE 	 ,
          E.5.1  Surface Condenser Design  	 ,
          E.5.2  Surface Condenser Cost  Estimation
                   Procedure. .... 	
D-5

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E-38
E-39

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     E.6
 PIPING AND DUCTING DESIGN AND COST ESTIMATION
 PROCEDURE 	
 E.6.1  Piping and Ducting Design Procedure. .
 E.6.2  Piping and Ducting Cost Estimation
          Procedure	
     E.7
APPENDIX

     F.I

     F.2
 REFERENCES FOR APPENDIX E 	
F - CALCULATION OF UNCONTROLLED EMISSION
    RATES AT SPECIFIC COST EFFECTIVENESSES
 PROCEDURE FOR INCINERATION DEVICES.
 F.I.I  General Procedure	
 PROCEDURES FOR CONDENSERS 	
 F.2.1  Styrene-in-Steam Emissions  .
 F.2.2  Styrene-in-Air Emissions .  ,
E-43
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E-54
E-59

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F-7
F-10
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                            LIST OF TABLES
                                                                 Page
2-1   End Uses of Polypropylene.	       2-4

2-2   Polypropylene (PP) Plants in Ozone
        Nonattainment Areas	       2-5

2-3   Characteristics of Vent Streams from the
        Polypropylene Continuous Liquid Phase Slurry
        Process	       2-9

2-4   Components of Polypropylene Vent Streams 	       2-10

2-5   High-Density Polyethylene (HOPE) Plants in Ozone
        Nonattainment Areas	  .       2-14

2-6   Characteristics of Vent Streams from the High Density
        Polyethylene Low-Pressure, Liquid Phase Slurry
        Process	       2-17

2-7   Polystyrene (PS) Plants in Ozone
        Nonattainment Areas	       2-20

2-8   Characteristics of Vent Streams from the Polystyrene
        Continuous Process 	 	       2-23

3-1   Flare Emissions Studies   ...... . . ......       3-11
4-1   Uncontrolled Emission Rates Versus Cost
        Effectiveness for Polypropylene Plants Based
        on Model Plant Parameters, by Process Section.  .  .       4-3

4-2   Uncontrolled Emission Rates Versus Cost
        Effectiveness for High-Density Polyethylene
        Plants Based on Model Plant Parameters,
        by Process Section	       4-4

4-3   Uncontrolled Emission Rates versus Cost
        Effectiveness for Polystyrene Plants
        Based on Model Plant Parameters,
        by Process Section	       4-5

4-4   Model Plant Environmental Analysis Based on
        Recommendations for RACT 	  .....       4-9
4-5   Additional Energy Required for Control
        with RACT in Polypropylene Plants	       4-11

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4-6   Additional Energy Required for Control
        with RACT in High-Density Polyethylene Plants.

4-7   Additional Energy Required for Control
        with RACT in Polystyrene Plants	

5-1   Cost Adjustments   	

5-2   Installation Cost Factors  	

5-3   Annualized Cost Factors for Polymers
        and Resins CTG	,, .  .
5-4   Polypropylene Model Plant Parameters
        and Emission Control Costs  	

5-5   High-Density Polyethylene Model Plant
        Parameters and Emission Control Costs	

5-6   Polystyrene Model Plant Parameters
        and Emission Control Costs  	

5-7   Cost Analysis for Polypropylene Model Plant.  .  .
5-8   Cost Analysis for Polypropylene Process
        Sections Across Process Lines	
5-9   Cost Analysis for Polypropylene Process
        Sections Within a Process Line 	
5-10  Cost Analysis for High-Density  Polyethylene   .  .

5-11  Cost Analysis for High-Density  Polyethylene
        Process Sections 	  .
5-12  Cost Analysis for Polystyrene Model  Plant.  .  .  .

5-13  Cost Analysis for Polystyrene Process
        Sections Within a Process Line	

5-14  Cost Effectiveness of RACT Applied to
        Continuous Streams in the Polymers and
        Resins  Industry, by Model Plant	

5-15  Cost Effectiveness of RACT Applied to
        Continuous Streams in the Polymers and
        Resins  Industry, by Process Section
        Across  Lines  	  	

5-16  Cost Effectiveness of RACT Applied to
        Continuous Streams in the Polymers and
        Resins  Industry, by Process Section
        Within  a Line	
D-l   Emission  Analyzers and  Instrumentation  Utilized
        for  Joint EPA/CMA Flare Testing	

D-2   Steam-Assisted  Flare Testing  Summary	
D-3   Summary of Thermal Incinerator  Emission Test
        Results	  .
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D-8
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D-4   Typical  Incinerator Parameters  for  ARCO  Polymers
        Emission Testing Based on  Data  From
        August 1981	       D-12

D-5   ARCO Polymers Incinerator Destruction Efficiencies
        for Each Set of Conditions ,	       D-16

D-6   Air Oxidation Unit Thermal  Incinerator Field Test
        Data		.  .       D-21

D-7   Destruction Efficiency Under Stated Conditions
        Based on Results of Union  Carbide
        Laboratory Tests .......<	*  .       D-34

D-8  Comparisons of Emission Test  Results
       for Union Carbide Lab Incinerator  and
       Rohm & Haas Field Incinerator	       D-38

E-l   Procedure to Design 98 Percent Efficient Elevated
        Steam-Assisted Smokeless Flares.  .........       E-4

E-2   Flare Budget Purchase Cost Estimates Provided by
        National Air Oil Burner, Inc.,
        in October 1982 Dollars.	       E-8

E-3   Capital and Annual Operating Cost Estimation
        Procedures for Steam-Assisted Smokeless Flares  .  .       E-12

E-4   Worksheet for Calculation of Waste Gas
        Characteristics	       E-16

E-5   Generalized Waste Gas Combustion  Calculations.  .  .  .       E-19

E-6   Procedure to Design Thermal  Incinerators
        Combusting Streams with Lower Heating
        Values Greater than 60 Btu/scf	       E-20
E-7   Capital and Annual Operating Cost Estimates for
        Retrofit Thermal Incinerators Without
        Heat Recovery	       E-25

E-8   Operating Parameters and Fuel Requirements
        of Catalytic Incinerator Systems	       E-29

E-9   Gas Parameters used for Estimating Capital and
        Operating Costs of Catalytic Incinerators	       E-31

E-10  Catalytic Incinerator Vendor Cost Data 	       E-33

E-ll  Capital and Operating Cost Estimation for
        Catalytic Incinerator Systems.  	       E-36
E-12  Procedure to Calculate Heat Load of a
        Condensation System for Styrene in Air	        E-40

E-13  Procedure to Calculate Heat Transfer Area of an
        Isothermal Condenser System	        E-44

E-14  Procedure to Calculate Heat Transfer Area of a
        Condensation System of Styrene in Air	        E-45

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E-15  Capital  and Annual  Operating Cost Estimates
        for a Retrofit 20 ft2 Condenser System for
        the Streams from the Continuous Polystyrene
        Model  Plant	       E-48
E-16  Capital  and 'Annual  Operating Cost Estimation
        Procedure for Condensers with Refrigeration. . .  .       E-49
E-17  Piping and Ducting Design Procedure	       E-55
E-18  Piping Components  	       E-56
E-19  Installed Piping Costs 	       E-57
E-20  Installed Ducting Cost Equations,
        1977 Dollars	  .       E-58
F-l   Initial Emission Characteristics and Control Costs
        for Calculation of Uncontrolled Emission Rates .  .       F-3
F-2   Summary of Annual Costs, $  	       F-5
F-3   Basic Minimum Costs at Various Flow Rates.  .....       F-6
F-4   Summary of Coefficients	       F-8
F-5   Summary of Cost Effective Flows  and Emission Rates,
        Polypropylene and High-Density Polyethylene
        Plants	       F-9
F-6   Control Costs for Styrene-in-Air-Emissions.  „ . .  . .      F-13
F-7   Exponents Used  for Condensers within Line  Analysis,
        $l,000/Mg  and $2,000/Mg	       F-16
F-8   Styrene-in-Air  Uncontrolled Emission Rates Equivalent
        to  $l,000/Mg, $2,000/Mg,  and $3,000/Mg  	       F-18

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                            LIST OF  FIGURES
2-1  Simplified Process Block Diagram for the
       Polypropylene Continuous,  Liquid Phase
       Slurry Process 	 . 	
2-2  Simplified Process Block Diagram for the
       High Density Polyethylene  Liquid Phase
       Slurry Process 	
2-3  Simplified Process Block Diagram for the
       Polystyrene Continuous Process. .	
3-1  Steam Assisted Elevated Flare System	
3-2  Steam Injection Flare Tip 	
3-3  Discrete Burner Thermal Incinerator 	  .  .
3-4  Distributed Burner Thermal  Incinerator	
3-5  Catalytic Incinerator 	
3-6  Condensation System 	 	
3-7  Two Stage Regenerative Adsorption System  	
3-8  Packed Tower for Gas Absorption	
D-l  Flare Sampling and Analysis  System  	
D-2  Schematic of Incineration System at ARCO
       Polypropylene Facility  	 ......
D-3  Incinerator Combustion Chamber  	
D-4  Petro-Tex Oxo Unit Incinerator  . 	
D-5  Off-gas Incinerator, Monsanto Co.,
       Chocolate Bayou Plant	. .  .  ,
D-6  Thermal Incinerator Stack Sampling System 	
E-l  Estimated Flare Purchase Cost for 40-ft. Height  .  .
E-2  Approximate Fluidic Seal Costs  	
E-3  Purchase Costs for Thermal  Incinerator Combustion
       Chambers	
E-4  Installed Capital Costs for Inlet Ducts, Waste Gas
     and Combustion Air Fans, and Stack for Thermal
     Incinerator Systems with No Heat Recovery 	
E-5  Installed Capital Costs for Catalytic Incinerators
       With and Without Heat Recovery	..'..,
E-6  Installed Capital Cost vs.  Condenser Area for
       Various Materials of Construction for a
       Complete Condenser Section	,
E-7  Installed Capital Costs vs.  Refrigeration Capacity
       at Various Coolant Temperatures for a Complete
       Condenser Section 	
                              x i
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                            1.0 , INTRODUCTION

     The Clean Air Act Amendments of 1977 require each State in which
there are areas in which the national ambient air quality standards
(NAAQS) are exceeded to adopt and submit revised State implementation
plans (SIP's) to EPA.  Revised SIP's were required to be submitted to
EPA by January 1, 1979.  States which were unable to demonstrate attainment
with the NAAQS for ozone by the statutory deadline of December 31, 1982,
could request extensions for attainment with the standard.  States
granted such an extension are required to submit a further revised SIP
by July 1, 1982.
     Section 172(a)(2) and (b)(3) of the Clean Air Act require that
nonattainment area SIP's include reasonably available control technology
(RACT) requirements for stationary sources.  As explained in the "General
Preamble for Proposed Rulemaking on Approval of State Implementation
Plan Revisions for Nonattainment Areas," (44 FR 20372, April 4, 1979)
for ozone SIP's, EPA permitted States to defer the adoption of RACT
regulations on a category of stationary sources of volatile organic
compounds (VOC) until after EPA published a control techniques guideline
(CT6) for that VOC source category.  See also 44 FR 53761 (September 17,
1979).  This delay allowed the States to make more technically sound
decisions regarding the application of RACT.
     Although CT6 documents review existing information and data concerning
the technology and cost of various control techniques to reduce emissions,
they are, of necessity, general in nature and do not fully account for
variations within a stationary source category.  Consequently, the
purpose of CTG documents is to provide State and local air pollution
control agencies with an initial information base for proceeding with
their own assessment of RACT for specific stationary sources.
                                 1-1

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                 2.0  PROCESS AND POLLUTANT EMISSIONS

2.1  INTRODUCTION
     The polymers and resins industry includes operations that convert
monomer or chemical intermediate materials obtained from the basic
petrochemical industry and the synthetic organic chemicals manufacturing
industry into polymer products.  Such products include plastic materials,
synthetic resins, synthetic rubbers, and organic fibers covered by
Standard Industrial Classification (SIC) codes 2821, 2822, 2823, and
2824.  The 1979 production of the major industry polymers was 16,052 Gg.
     Thirty-six percent of this total production of the industry is from
the manufacture of high-density polyethylene, polypropylene, and polystyrene.
In addition, the manufacture of these three polymers is estimated to
account for 56 percent of the total estimated industry process emissions
of 86.2 Gg/yr of volatile organic compounds (VOC).
     This chapter describes the manufacturing processes for each of
these three polymers under consideration and the associated process VOC
emissions.  In general, the manufacture of these polymers may be
considered as a five step operation: (1) raw materials storage and
preparation, (2) polymerization reaction, (3) material recovery, (4)
product finishing, and (5) product storage.  The equipment used in each
process step may have associated process emissions.  The relationship
between process section (that is, the group of equipment used in
the performance of one of the five basic process steps) and process
emissions is shown in the tables identifying vent stream characteristics
for each polymer type.
     Fabrication, blending, or formation of resin materials are not
included in the process descriptions, nor are emissions from these
operations quantified.  Fugitive and storage emissions from these
processes are described in other CTG documents, "Control of Volatile
Organic Fugitive Emissions from Synthetic Organic Chemical and Polymers
                                  2-1

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and Resins Manufacturing Equipment" and "Control  of Volatile Organic
Emissions from Volatile Organic Liquid Storage in Floating and  Fixed
Roof Tanks" and hence, they are not discussed here.
     The model plants in this chapter represent most of existing processes
in the ozone nonattainment areas for each particular resin.  The uncontrolled
emission factors can be used as a basis for the verification of VOC
emissions developed from emission source tests, plant site visits,
permit applications, etc.  These emission factors should not be applied
in cases where site-specific data are available, but rather, in instances
where specific plant information is lacking or highly suspect.   States
may choose to analyze, or EPA may subsequently analyze, other processes
not represented by the model plants, such as the relatively new gas phase
processes of polypropylene and polyethylene or the less common  liquid phase
solution process of high-density polyethylene production.
2.2  POLYPROPYLENE
2.2.1  General Industry Description
     Manufacture of polypropylene, on  a commercial scale, started in the
1950's when stereospecific catalysts were discovered.  Polypropylene is
a  high-molecular weight thermoplastic  crystalline  polymer of propylene.
The general formula for polypropylene  is as  follows:
                                                                          1
                                                                          I
                .  .  .  CH2  - CH  -  CH2  -  CH - CH2 - CH  -  .  .  .
                            i           i           i
                           CH3       CH3         CH3
     The polymer  is lightweight, water-  and  chemical-resistant,  somewhat
 rigid,  and easy to  process.   It  exists in three  different  forms  depending
 on the  geometric  arrangement of  the  methyl groups:  (1)  isotactic  - with
 all methyl  groups  aligned on the same side of the  chain  as  shown above,
 (2)  syndiotactic  -  with the  methyl  groups  alternating,  and  (3)  atactic  -
 all  other forms  in  which  the methyl  groups are randomly aligned on
 either  side of the  chain.  Typically, commercial  polypropylene  consists
                                                                          i
 principally of crystalline material  (isotactic), with  only a small amount
 of amorphous  material  (atactic).!                                         i
      Consumer products from polypropylene  can be formed in  many ways,
 including solid molding,  extrusion,  rotational  molding,  powder  watering,
 thermoforming, foam molding, and fiber orientation.2
                                   2-2

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     Polypropylene resins are supplied in many grades for a variety  of
uses.  Apart from major distinctions between homopolymer, intermediate-impact
co-polymer, and high-impact co-polymer, the grades may also differ in
specific formulations.  Different grades of polypropylene lend themselves
to use in different applications.  Molded applications include bottles
for syrups and foods, caps, auto parts, appliance parts,  toys, housewares,
and furniture components.  Fibers and filaments are used  in carpets,
rugs, and cordage.  Film uses include packaging for cigarettes, records,
and housewares.  Extrusion products include pipes, profiles, wires and
cable coatings, and corrugated packing sheets.3
     Injection molding accounts for 41 percent of polypropylene use;
fibers and filaments account for 31 percent; and other forms account for
28 percent.3  In terms of end uses, major sectors are shown in Table 2-1.
     Production of polypropylene has grown from 981 Gg in 1973 to 1,743-Gg
in 1979, a 10.1 percent annual growth rate.  C.H. Kline projects a
9.0 percent growth rate for polypropylene from 1978 to 1983,4 and SRI
International projects an 8 percent growth rate from 1977 to 1982.5
Currently, 24 plants produce polypropylene in the United  States.6  The
existing polypropylene plants known to be in the current  ozone nonattainment
areas are listed in Table 2-2.
2.2.2  Model Plant
     The continuous slurry process for manufacture of polypropylene is
the most widely used process commercially.  Based on data from 10 existing
plants located in nonattainment areas, a model plant capacity of 141 Gg/yr
was selected.
     The polypropylene resins, characterized by having a  controlled
content of isotactic material, are obtained through coordination polym-
erization, employing a heterogeneous Ziegler-Natta type catalyst system,
which typically is a combination of titanium tetrachloride and aluminum
alkyls.  More recent process technology, which uses a high-yield catalyst
with improved activity, requires much less catalyst than  the conventional
process.  With this high-yield process, the catalyst is left in the
product.  This technology results in fewer processing steps and, thus,
less emissions.  This new process is incorporated in the  model plant by
exclusion of several processing units, and is consistent  with a proportional
reduction in the total emission factor.
                                  2-3

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                  Table 2-1.  END USES OF POLYPROPYLENE
       Sector
 Weight Percent
Polypropylene Use
Consumer/Institutional
Furniture/Furnishings
Packaging
Transportation
Electrical/Electronics
Other
        19
        18
        16
        12
         7
        28
                                  2-4

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   Table  2-2.   POLYPROPYLENE  (PP)  PLANTS  IN  OZONE  NONATTAINMENT AREAS3

Company
ARCO Polymers, Inc.
Amoco Chem. Corp.
Exxon Chem. Co.
Gulf Oil
Hercules, Inc.
Northern Petrochem.
Co.
USS Novamont Corp.
Phillips Petro. Co.
Rexene Polyolefins
Co.
Shell Chem. Co.
Soltex Polymer Corp.
Location
Deer Park, TX
Chocolate Bayou, TX
Baytdwn, TX
Cedar Bayou, TX
Bay town, TX
Lake Charles, LA
Morris, IL
La Porte, TX
Pasadena, TX
Odessa, TX
Bay port, TX
Norco9 LA
Wood bury, NJ
Deer Park, TX
Status5
NAR
NANR
NAR
NAR
NAR
NANR
NANR
NAR
NAR
NANR.
NAR
NANR
NAR
NAR
Capacity
(Gg/yr)
181
125
250
181
272
376
91
159
82
23-46
136
136
91
 This list is illustrative only.   Since the nonattainment status of
 areas changes from time to time, this is n9t intended to be a definitive
 list of plants that will be affected by this guideline document.

 Ozone nonattainment area not requesting extension (NANR).
 Ozone nonattainment area requesting extension (NAR).


SOURCES:  SRI International, 1980 Directory of Chemical Producers,
          United States.

          U.S. EPA study by Pullman-Kellogg Co., plant listing.

          The BNA Environmental Reporter AQCR Listing.  §121 (through
          March 12, 1981).
                                 2-5

-------
     2.2.2.1  Process Description.  The continuous slurry processes,
conventional and high-yield, are represented in Figure 2-1.   Reactor
feed materials consist mainly of monomer propylene, comonomer ethylene,
monomer impurities propane and ethane, hexane, and a stereospecific
catalyst.  Hexane is used as a process diluent and acts as a heat transfer
agent and polymer suspending medium.  The catalyst is usually manufactured
on site to consistently maintain the required catalyst activity.   It  is
mixed with necessary solvents and metered accurately into the polymerization
reactor along with other reactants.  Process diluent is also used in
catalyst preparation and spent diluent is sent to the diluent recovery
section for reuse.
     The reactor is a continuously stirred jacketed vessel or a loop
reactor.  During reaction, a portion of the polymer/monomer/diluent
mixture is continuously drawn from the reactor to a flash tank in which
the unreacted propylene and propane are separated, and recovered by
condensation.
     Slurry from the flash tank is then fed to the deactivation/decanting
section for washing with an alcohol-water solution to remove most of  the
catalyst residues.  The diluent/crude product slurry is lighter than  the
alcohol-water solution and the two phases are separated by decantation.
The alcohol-water phase is distilled to recover alcohol; whereas, the
diluent/crude product phase which is in the form of a slurry is stripped
to remove part of the diluent.  The product slurry is then sent to a
slurry vacuum filter system in which isotactic polymer product solids are
separated from the diluent.  The atactic polymer remains dissolved in
the diluent.  The isotactic product goes through a product dryer, then is
extruded, pelletized, and sent to product storage.
     In the methanol recovery section, the crude methanol streams are
                                               ,.                  '       I '
refined and recycled, and the bottom streams, containing catalysts
metals are sent to the plant waste-water treatment facility.
     The atactic-diluent solution is fed to the by-product (atactic)  and
diluent separation unit in which the diluent is purified and dried for
recycle, and the atactic solids are recovered or burned in incinerators.
     In the high-yield slurry process, the catalyst is left in the
product so deactivation/decanting and alcohol recovery sections are
unnecessary.  Along with this, one of the major emission streams is also
                                  2-6

-------
eliminated.  Figure 2-1 indicates the units that should be excluded in
this process.
     In addition to the use of high-yield catalysts,  other process
variations may occur.  Mixtures of aliphatic hydrocarbons may replace
hexane as the process diluent, and isopropyl alcohol  may replace methanol
as the catalyst deactivation agent.  Also polymer dryers may vary with
the facility, but a fluid bed dryer with hot nitrogen or air is the most
common.  Nitrogen drying has safety advantages and the nitrogen can be
recycled resulting in lower costs and emissions.  Other types of product
dryers and different operating pressures may result in a much higher VOC
emission rate.  Except for high-yield catalyst, and the product dryer
type and operating pressure, these other process variations are minor
and should have little effect on the process VOC emissions.
     2.2.2.2  VOC Sources.  The offgas stream characteristics for polypropylene
manufacturer are shown in Table 2-3.  The combined process VOC emission
factor for the conventional slurry process is 36.7 kg VOC/1,000 kg
product.  For the high-yield slurry process, Streams C and D are not
present; therefore, the combined process VOC emission factor for this
process is about 23.4 kg VOC/1,000 kg product.  Most of the emission
streams are continuous and consist mainly of propylene, ethylene, propane,
and a small amount of process diluent.  Properties of these compounds
are summarized in Table 2-4.  The temperature of the streams varies from
ambient to 1040C, and the pressure is about atmospheric.  Each of the
major VOC-containihg streams are indicated on Figure 2-1 and are described
below:
     1.  Stream A:  Catalyst Preparation Vents - This vent continuously
releases process diluent that is used in preparation of the catalyst.
     2.  Stream B:  Combined Polymerization Reactor Vents - These emissions
are from vents of reactors from all process trains.  This is a continuous
stream venting organic process offgas, consisting mainly of C3 (propylene
monomer and other hydrocarbons with three carbon atoms such as propane)
and process diluent, which could be hexane or a mixture of aliphatic
hydrocarbons with 10-12 carbon atoms.
     3.  Streams C & D:  Decanter and Neutralizer Vents - These vents
are part of the alcohol recovery section.  This is usually the largest
VOC source in the process and consists of methanol or isopropyl alcohol,
                                  2-8

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          Table 2-4.   COMPONENTS OF POLYPROPYLENE VENT STREAMS
Propylene (monomer)
Propane (monomer impurity)
n-Hexane (diluent)
Methanol or Isopropanol
  (washing alcohol)
Ethylene (comonomer).
Co-Ce Hydrocarbons (might include
  ethylene, propylene, and propane)
C10 H.C. (A mixture of aliphatic
  Hydrocarbons with 10-12 carbon
  atoms.)
                                         MU = 42.06,  2186 Btu/cu ft
                                         MW = 44.09,  2385 Btu/cu ft
                                         MW = 86.17,  4412 Btu/cu ft
                                         MW = 32.04 or 60.02

                                         MW = 28.05,  1513 Btu/cu ft

                                         MW = 50 (Avg)
                                         MW = 144.0
All of these compounds are usually diluted in gases like:
Air                                      MW = 29.0
Nitrogen                                 MW = 28.0
Hydrogen                                 MW = 2.0, 275 Btu/cu ft
                                  2-10

-------
in addition to €3 and process diluent.  The stream is continuous and
exists in most of the existing polypropylene plants.  The process using
a high-yield catalyst does not require these vents, and the reduction in
total emission factor is significant.
     4.  Stream E:  Slurry Vacuum Filter System Vents - This stream is
from the system which separates the atactic and isotactic polymer.  It
is one of the largest VOC emission streams venting process diluent and
alcohol remaining in the polymer.  It is a continuous stream at atmospheric
pressure and exists in both the conventional and high-yield slurry
process plants.
     5.  Stream F:  Diluent Separation and Recovery - This stream originates
from the by-product and diluent recovery section and can be the second
largest VOC emission stream in the entire process.  The diluent recovery
section which consists of an evaporator, an extractor and distillation
units is part of all processes and emits process diluents and alcohol
vapors.
     6.  Stream G:  Dryer Vents - This vent emits hydrocarbons diluted
in air or nitrogen at a relatively high temperature (104°C) and atmospheric
pressure.  The emissions consists of vapor of hexane, methanol, and
propane.
     7.  Stream H:  Extrusion/Pelletizing Vent - This vent can continuously
emit significant quantities of hydrocarbon that may remain in the polypropylene
powder as it exits the dryer and enters the extruder feed chute.  At
this point, the powder is in equilibrium with a vapor that can contain
up to 25 percent hydrocarbon by weight.  As a result of heating and
compression in the extruder, there is some VOC loss through the extruder/
pelletizing section and futher losses from the powder/pellet transfer
system downstream from the product dryer since the transfer medium acts
as a stripping gas.
     The stream properties and VOC concentrations of Streams A to H can
vary depending on process conditions.  The variation generally depends
on the product grade or type being manufactured and other variables such
as temperature, pressure, catalyst concentration or activity, and the
amount of hydrogen used for molecular weight control.  The concentration
and the magnitude of each stream is, of course, hightest under start-up
or shutdown conditions because of process conditions away from equilibrium.
                                  2-11

-------
     2.2.2.3  Control Systems.  No controls are routinely  applied  for
VOC control of these continuous sources.   The polymerization  reactors
and the atactic separation units, however, are generally provided  with
emergency relief valves leading to a flare for safety purposes  in  the
case of upsets.  These emergency vents usually pass through  knock-out
drums to separate entrained liquid and polymer particles before the
vapors are piped to the flare.  Also, in  the production steps,  the
concentrated atactic polymer stream from  the slurry vacuum filter  system
1s piped to a vessel and its liquid content is removed by  evaporation.
The solid amorphous atactic polypropylene is left behind and  is then
either burned in incinerators or is packed and sold as a by-product for
paper coating and other applications.  For some producers, the  atactic
polymer is incinerated, liquid and gaseous waste streams from the  process
may also be burned in the same device.
2.3  HIGH-DENSITY POLYETHYLENE
2.3.1  General Industry Description
     High-density polyethylene (HOPE) resins are linear thermoplastic
polymers of ethylene with densities higher than 0.94 g/cm3.   HOPE  resins
                                                                      i
are typically produced by a low-pressure  process in which  organic  solvents
are used; the solid catalyst is in suspension; and the polymer  forms a
slurry (e.g., the processes originated by Phillips Petroleum Company and
Solvay and Cie, SA).  Although there are various solvent processes used,
the variations do not affect emissions except with respect to the  solvent
                                                                      i
recovery methods used.
     HOPE is a highly (>90 percent) crystalline polymer containing less
than one side chain per 200 carbon atoms  in the main chain.   The typical
density range is 0.95-0.97 g/cm3.?  it is strong, water- and chemical-
resistant, and can  be easily processed.  It is one of the largest  volume
plastics produced in the U.S. and in the world.  It is extruded into
film sheets, pipe or profiles, coated, injection molded, blow molded,
rotationally molded, foamed, or formed in other ways.2
     HDPE's primary application is blow molded bottles for bleaches,
liquid detergents,  milk, and other fluids.  Other blow molded forms for
which HDPE's are used include automotive gas tanks;, drums, and  carboys.
HDPE's also are used for injection molded forms including material
handling pallets, stadium seats, trash cans, and auto parts.  Film is
                                  2-12

-------
used in making shopping bags.  Forty percent of all  HOPE  is blow molded;
another 22 percent is injection molded.   Film and sheet combined account
for only six percent of HOPE use.  Other uses account for 32 percent.
End use sectors for HOPE include packaging (45 percent),  consumer/insti-
tutional (11 percent), building and construction (9  percent), and other
sectors (35 percent).3
     From 1973 to 1979, production of HOPE grew from 1,196 Gg to 2*273 Gg,
a growth rate of 11.3 percent.  C.H. Kline projects  growth at 7.0 percent
for 1978 to 1983.4  SRI International projected growth from 1976 to 1980
at 10 percent.5
2.3.2  Model Plant
     The Phillips particle form process serves as the basis for this
model plant, but it is intended to represent all other liquid phase
slurry processes.
     This model plan specifically includes an unreacted monomer recycling
system.  There are other similar liquid-phase processes that do not use
such systems and have larger emissions.  The plant capacity for the
model HOPE plant is 214 Gg/yr.  this is based on plants located in
nonattainment areas.  The existing HOPE plants known to be in the current
ozone nonattainment areas are listed in Table 2-5.
     2.3.2.1  Process Description.  Referring to the schematic for this
process, Figure 2-2, the feed section includes catalyst purification and
activation.  The prepared catalyst is then fed to the reactor continuously
by being slurried in a stream of process diluent (pentane or isobutane).
Ethylene monomer and comonomer  (1-butene or hexene), after purification,
are also fed to the reactor where polymerization takes place in process
solvent.  The reactor, for the particle-form process, is  usually a
closed loop pipe reactor.  The product HOPE is separated from unreacted
monomer and diluent by flashing from a low pressure to a vacuum and by
steam stripping.  The wet polymer solids are dewatered in a centrifuge
and then dried in a closed-loop nitrogen or air-fluidized drying system
prior to extrusion.
     The unreacted monomer and diluent vapors are sent through a diluent
recovery unit where most of the diluent is separated and recycled back
to the reactor.  The rest of the stream is then sent to the ethylene
                                  2-13

-------
     Table  2-5.   HIGH-DENSITY  POLYETHYLENE  (HOPE)  PLANTS  IN  OZONE
                          NONATTAINMENT  AREAS3

Company
Allied Chem. Corp.
ARCO Polymers, Inc.
Cities Service Co.
Dow Chemical
Amoco Chem. Corp.
E.I. Du Pont de
Nemours & Co. Inc.
Gulf Oil Corp.
Hercules, Inc.
Nat1! Petrochem.
Corp.
Phillips Petro. Co.
Soltex Polymer Corp.
UCC
Location
Baton Rouge, LA
Port Arthur, TX
Texas City, TX
Freeport, TX
Plaquemine, LA
Chocolate Bayou, TX
Orange, TX
Victoria, TX
Orange, TX
Lake Charles, LA
La Porte, TX
Pasadena, TX
Deer Park, TX
Port Lavaca, TX
Status13
NANR
NANR
NANR
NANR
NANR
NANR
NANR
NANR
NANR
NANR
NAR
NAR
NAR
NANR
Capacity
(Gg/yr)
272
!
147
82
136 :
136
159
104 !
}02
1
200
7
227 '
420
270
181
aThis list is illustrative only.  Since the attainment status of
 areas change from time to time, this is not intended to be a definitive
 list of plants that will be affected by this guideline document.
^Ozone nonattainment area not requesting extension (NANR).
 Ozone nonattainment area requesting extension (NAR).

SOURCES:    SRI International, 1980 Directory of Chemical Producers,
            United States.

            U.S. EPA study by Pullman-Kellogg Co., plant listing.

            The BNA Environmental Reporter AQCR Listing.  §121 (through
            March 12, 1981).
                                 2-14

-------
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recovery unit where ethylene is recovered and sent to recycle ethylene
treaters and back to the reactor.
     2.3.2.2  VOC Sources.  All the process streams, except the feed
preparation stream, in HOPE manufacture are continuous, and they consist
mainly of ethylene and process solvent diluted in nitrogen or air.
Most of the streams are at ambient temperature.  An ethylene safety
flare is always a part of each system, and some plants may use it for
VOC emission control.  Since this particular model plant incorporates
ethylene recycle, it has relatively small emissions, but plants which
vent unreacted monomer and use air-fluidized dryers have substantially
higher VOC emissions.  The major VOC source is the flash tank where an
unreacted monomer stream (about 50 percent VOC) is released.  HOPE
manufacturers often send this stream to a boiler to recover the heat
                                                             , ,           i
content.  Table 2-6 shows the vent stream-characteristics for the VOC    ]
sources:  these sources are described below:
                                                                         i
     1.  Stream A:  Feed Preparation - This is an intermittent stream
consisting mostly of ethylene.  Assumed to vent 12 times a year, it's
sources are drying, dehydrating and other feed purification operations.
     2.  Stream B:  Dryer - Dryer emissions are continuous and have low
VOC concentrations.  Closed-loop drying systems have very low emissions
of process solvent in nitrogen.  Air-fluidized dryers have significantly
higher emissions.
     3.  Stream C:  Continuous Mixer - This is another low VOC emission
stream coming from a mixer which mixes polymer with anti-oxidants.   It
is continuous and releases process solvent that is still left in the
polymer along with a large quantity of nitrogen.  Usually this stream is
emitted to the atmosphere.
     4.  Stream D:  Recycle Treaters - This is a  semi-continuous VOC
emission stream containing about 80 weight percent VOC.  Currently this
stream is usually flared.  Treaters consist of vessels containing such
materials as adsorbents, dessicants, and molecular sieves which remove
water and other impurities in the recycle ethylene stream.  Emissions
occur when the vessels are purged during  regeneration of the adsorber
beds.  This stream is considered a continuous stream.  The stream flows
continuously for about 20 out of 24 hours.
                                  2-16
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     2.3.2.3  Control  Systems.   As noted,  like the  other polyolefin
processes, the HOPE process generally has  a flare as  a part  of  the
system for safety reasons.  A complete line of safety relief devices
leading to the flare are commonly provided to avoid accidents as  a
result of equipment overpressurization or  malfunction.
2.4  POLYSTYRENE
2.4.1  General Industry Description
     Polystyrene offers a combination of excellent  physical  properties
and processibility at a relatively low price for thermoplastic  materials.
It is crystal clear and has colorability,  rigidity, good electrical
properties, thermal stability, and high-flexural and tensile strengths.
Polystyrene products are used in molded forms, extrusions, liquid solutions,
adhesives, coatings, and foams.  The family of polymerized co-polymers
                                                       i        •         i
from styrene monomer and its modifications ranked third among all plastics
in consumption within the United States.8
     Molded uses include toys, autoparts, housewares, kitchen items,
appliances, wall tiles, refrigerated food containers, radio and television
housings, small appliance housing, furniture, packages, and building
components such as shutters.  Extruded sheets also are used in  packaging,
appliance, boats, luggage, and disposable plates.  F:oamed styrene is  a
good insulator and is used in construction, packaging, boats, housewares,
toys, and hot/cold insulated drink cups.2  Fifty percent of all styrene
                                                      1 ' '     •',"''     i
is molded.  Extrusion accounts for 33 percent.  Other forms make up
17 percent.
     Of end use sectors,  packaging makes  up 35  percent, consumer/
institutional - 22 percent, building and  construction - 10 percent,
electrical  and electronic - 10 percent, and other  sectors - 23 percent.3
     Production of styrene has grown from 1,507 Gg in 1973 to 1,817 Gg
in 1978,  a  3.2 percent  growth  rate.  C.H.  Kline projects a 6.0 percent
growth rate for 1978-19834 while  SRI  International projects a 4.9 percent
growth rate for 1979-1982.5
     Styrene  polymerizes  readily  with the addition of either heat or
catalyst  like benzoyl peroxide or ditertiary  butylperbenzoate.   Styrene
will homopolymerize in  the presence  of  inert  materials  and co-polymerize
with a variety of monomers.  Pure polystyrene has  the following  structure:
                                   2-18
-------
                 H
                 C
- CH2 - C - CH2 - C - CH2 - C -
       666
     Although polymers with molecular weights in the millions can be
made, those most useful for molding have molecular weights of about
125,000; while those used in the surface coating industry average about
35,000.
2.4.2  Model Plant
     A continuous process for the manufacture of polystyrene was chosen
for developing the model plant primarily because of its significant VOC
emissions.  Mass (bulk) polymerization was used as a basis for developing
the flow diagram.  However, the model plant represents all liquid phase
continuous processes.  In the case of suspension polymerization, because
polymerization takes place in water, dewatering, washing, centrifuge and
dryer sections are required.  These sections usually are not sources of
VOC emissions.  The model plant capacity is 73.5 Gg/yr.  This capacity
represents an average of capacities from polystyrene plants using batch
or continuous processes in ozone nonattainment areas.  The existing
polystyrene plants in ozone nonattainment areas are listed in Table 2-7.
The list includes both continuous and batch-type processes; when the
process type is unknown the process comment is left blank.  The plants
with unknown process type are included for completeness of the list.
Only the continuous processes are covered by RACT.
     2.4.2.1  Process Description.  This description is for a
fully continuous, thermal co-polymerization process for the manufacture
of pelletized polystyrene resin from styrene monomer and polybutadiene.
Several grades of crystal and impact polystyrene are produced by this
process.  The continuous process is represented in Figure 2-3.
     Styrene, polybutadiene, mineral oil, and small amounts of recycle
polystyrene, anti-oxidants and other additives are introduced into the
feed dissolver tank in proportions that vary according to the grade of
resin being produced.  Blended feed is pumped on a continuous basis to
the  reactor where the  feed is thermally polymerized to polystyrene.  The
polymer melt, containing some unreacted styrene monomer and by-products
                                  2-19
-------
Table  2-7.    POLYSTYRENE (PS) PLANTS  IN OZONE NONATTAINMENT  AREAS'
Company
A.E. Plastfk Pak Co., Inc.
Am. Hoechst Corp.
Amoco Cheralcal Corp.
ARCO Polymers, Inc.
BASF Wyandotte Corp.
Carl Gordon, Ind., Inc.
Cosden Oil & Chemical Co.
Crest Container Corp.
Dart Ind., Inc.
Dow Chaiical Corp.
Gulf 011 Chemical Co.
Mobil Chemical Co.
Monsanto
Polysar Resins, Inc.
Richardson Company
Shell Chemical Co.
Sterling Plastics Corp.
Location
City of Industry, CA
Cheasapeake, VA
Leonvfnster, HA
Jollet, IL
Torrance, CA
WHlow Springs, IL
Honaca, PA
James burg, NJ
South Brunswick, NJ
Owens bo ro, KY
Oxford, HA
Worchester, HA
Windsor, NJ
Calumet City, IL
Saginaw, TX
Fort Worth, TX
Bayport, TX
Allyns Pt., CT
Midland, MI
Torrance, CA
HaMetta, OH
Channel view, TX
Holyoke, MA
Joilet, IL
Santa Ana, CA
Addyston, OH
Decatur, AL
Long Beach, CA
Springfield, HA
Copley, OH
Leorainster, MA
Channel view, TX
Belpre, OH
Windsor, NJ
Statusb
NAR
NANR
NANR
NANR
NAR
NAR
NAR
NAR
NAR
NAR )
NAR >
NAR )
NAR
NAR
NANR
NANR
NANR
NAR
NANR
NAR
NANR
NAR
NAR
NANR
NANR
NAR
NANR
NAR
NAR
NANR
NAR
NANR
NANR
NAR
Capacity
(Gg/yr)
16
91
54
136
16
41
238
136
50
68
54
120
14
3.6
68
82
100
91
	 102 	
18
45
20
34
136
45
23
136
82
52
-
141
13.6-
54.4
Process
Comment0
-
-
Continuous
Batch
Batch
-
Batch
Batch
Batch
Continuous
Batch
- •
Continuous
Continuous
Continuous
Continuous
Continuous
Continuous
Continuous
Continuous
-
Continuous
Continuous
 aTh1s  11st is illustrative only.  Since the attainment status of areas  change
  from  time to time, this is not intended to be a definitive 11st of plants that will
  be affected by this guideline document.
  Ozone nonattainment area not requesting extension  (NANR).
  Ozone nonattalnment area requesting extension (NAR).
 C0nly  continuous processes are covered by RACT.


         SOURCES;    SRI International, 1980 Directory of Chemical Producers,
                    United States.

                    U.S. EPA study by Pullman-Kellogg Co., plant listing.

                    The BNA Environmental Reporter  AQCR Listing.  §121  (through March 12,
                    1981).
                                              2-20
                                                                                           .  i	
-------
                                                              CO
                                                              GO
                                                              
2-21
-------
is pumped to a vacuum devolatilizer where most of the monomer and by-
products are separated, condensed and sent to a styrene recovery unit.
Vapors from the styrene condenser are vented through a vacuum system.
     Molten polystyrene from the bottom of the devolatilizer is pumped
through a stranding die-plate into a cold water bath.  The cooled strands
are pelletized and sent to product storage.
     In the styrene recovery unit, crude styrene monomer is separated in
                                                 . .   .;!   i     '-         I    J
a distillation column.  The styrene vapor overhead from the tower is
condensed and recycled to the feed dissolver tank.   Noncondensibles are
vented through a vacuum system.  Heavies from the bottom of the column
can be used as a fuel supplement.
     2.4.2.2  VOC Sources.  Table 2-8 shows the vent stream characteristics
for the continuous polystyrene process.  All VOC emission streams from
the process are continuous.  Industry's experience with continuous
polystyrene plants indicate a wide range of emission rates from plant to
plant.  Steam present in Streams B and C reflects the use of a steam jet
ejector in the vacuum system used;  air reflects the use of vacuum
pumps.
     1.   Stream A:  Feed Dissolver - This vent emits mostly styrene.
The VOC emission results from washing losses.  Currently, the styrene is
emitted to the atmosphere.
     2.   Stream B:  Styrene Condenser Vent - Consists of unreacted
styrene separated from the polystyrene in a vacuum devolatilizer.  The
                                                                      j
stream can be exhausted through a vacuum system (e.g., steam jet ejector)
to atmosphere.  This is the largest VOC source.  When vacuum pumps are
used and followed by refrigerated brine condenser,, the emissions can be
lower.
     3.   Stream C:  Styrene Recovery Unit Condenser Vent - This stream
contains the noncondensible components separated in the styrene recovery
tower and is vented through a steam jet ejector or vacuum pump.
     4.   Stream D:  Extruder Quench Vent - This stream consists of
steam and a trace of styrene vapor.  The stream is usually vented through
a forced-draft hood and passed through demister-pad or electrostatic
precipitator before venting to the atmosphere.
     2.4.2.3  Control Systems.  No routine control is applied to continuous
                                                                      j
processes other than normal condensation operations.  One unique system,
                                  2-22
-------







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however, of vapor condensing/recovery is used where each process vessel
1s equipped with rupture discs having the respective pressure relief
settings.  When any of these process vessels are overpressured, the
                                                                        !
vapors relieve to the vapor condensing/recovery system.  By flashing
action and by condensation, most of the vapors are condensed, recovered,
and reused in the process.  This system also results in a single emission
point in the entire process.  Unlike the polyolefins; processes, no
flares are used as control devices.
                                  2-24
-------
2.5  REFERENCE FOR CHAPTER 2

1.   Letter from Matey, J.S., Chemical Manufacturers Association, to
     Vincent, E.J., EPA.  October 19, 1981.  p. 2.  CMA comments on EPA
     model plant emission factors for polymers/resins manufacture.

2.   Kline, C.H. and Company.  The Kline Guide to the Plastics Industry.
     Fairfield, New Jersey. 1978.

3.   SRI International.  Facts and Figures of the Plastics Industry.
     New York, New York. 1978.

4.   Kline, C.H. and Company.  Plastics and Resins:  Forecast to 1983.
5.

6.



7.
Chemical and Engineering News reprint. 1979.

SRI International.  Chemical  Economics Handbook.

Click, C.N. and O.K. Webber,  Polymer Industry Ranking by VOC Emissions
Reduction that Would Occur from New Source Performance Standards.
Pullman-Kellogg Company, EPA  Contract No. 68-02-2619.  1979.  p.  177-178.

Billmeyers, F.W.  Textbook of Polymer Science.  New York, Wiley
Interscience, 1971.  p. 379-386.
8.   SRI International.  The Story of the Plastics Industry.  1977.
                               2-25
-------
-------
                      3.0  EMISSION CONTROL TECHNIQUES

     Volatile organic compounds (VOC), used as solvents and key raw
materials in the manufacture of polymers and resins, are emitted to the
atmosphere from a variety of process equipment.  Process VOC emissions
can be reduced either by installing emission control devices or by
reducing the VOC in the vent streams by a process modification such as
recovery of monomer or solvent.  This chapter describes emission control
techniques that may be used to reduce process emissions from the polymers
and resins industry.
     Process emissions from the manufacture of polymers and resins are
diverse in both composition and flow.  Streams contain a wide range of
VOC concentrations, i.e., less than 1 percent to essentially 100 percent,
but most are of high concentration.  Some streams are continuous, while
others are intermittent.  Process emissions also differ in temperature,
pressure, heating value, and miscibility.  These factors are extremely
important in the selection and design of VOC emission control equipment.
     Due to this diversity, different control techniques may be appropriate
for different vent streams.  The control techniques may be characterized
by two broad categories:  combustion techniques and recovery techniques.
Combustion techniques such as flares and incinerators are applicable to
a variety of VOC streams.  Recovery techniques such as condensation,
absorption, and adsorption, are effective for some select vent streams.
Economic incentives may encourage the use of either type of VOC control,
since certain combustion configurations may permit heat recovery, and
recovery techniques permit the conservation and reuse of valuable materials.
The selection of a control system for a particular application is based
primarily on considerations of technical feasibility and process economics.
     The most common control techniques form the basis for this chapter.
Basic design considerations for flares, thermal and catalytic incinerators,
                                 3-1
-------
industrial boilers, condensers, absorbers, and adsorbers, are briefly
described.  The conditions affecting the VOC removal efficiency of each
type of device and its applicability for use in the polymers and resins
Industry are examined.  Emphasis has been given to flares, thermal
incinerators, and condensers because of their wide applicability to a
variety of VOC streams.  Combustion techniques are discussed in Section 3.1
and recovery techniques in Section 3.2.
3.1  CONTROL BY COMBUSTION TECHNIQUES
     The four major combustion devices that are or can be used to control
VOC emissions from the polymers and resins industry are:  flares, thermal
or catalytic incinerators, and boilers.  Flares are the most widely used
control devices at polyethylene and polypropylene manufacturing plants.
Incinerators and boilers are also used, to a lesser extent, to control
continuous vent streams.  Although these control devices are founded
                                                                         .
upon basic combustion principles, their operating characteristics are
very different.  While flares can handle both continuous and intermittent
streams, neither boilers nor incinerators can effectively handle large
volume intermittent streams.  This section discusses the general principles
of combustion, and then the design and operation, VOC destruction efficiency,
                                                                         I
and applicability of these four combustion devices at polymers and
resins manufacturing plants.
     Combustion is a rapid oxidation process, exothermic in nature,
which results in the destruction of VOC by converting it to carbon
dioxide and water.  Poor or incomplete combustion results in the production
of other organic compounds including carbon monoxide,.  The chemical
reaction sequence which takes place in the destruction of VOC by combustion
is a complicated process.  It involves a series of reactions that produce
free radicals, partial oxidation products, and  final combustion products.
Several intermediate products may be created before the oxidation process
is completed.  However, most of the intermediate products have a very
short life and, for engineering purposes, complete destruction of the
VOC is the principal concern.
     Destruction efficiency is a function of temperature, turbulence,
and residence time.  Chemicals vary in the magnitudes of these parameters
                                  3-2
-------
that they require for complete combustion.  An effective combustion
technique must provide:2
     1.  Intimate mixing of combustible material (VOC) and the oxidizer
(air),
     2.  Sufficient temperature to ignite the VOC/air mixture and complete
its combustion,
     3.  Required residence time for combustion to be completed, and
     4.  Admission of sufficient air (more than the stoichiometric
amount) to oxidize the VOC completely.
3.1.1  Flares
     Flaring is an open combustion process in which the oxygen required
for combustion is provided by the air around the flame.  Good combustion
in a flare is governed by flame temperature, residence time of components
in the combustion zone, turbulent mixing of components to complete the
oxidation reaction, and oxygen for free radical formation.
     There are two types of flares:  ground level flares and elevated
flares.  Kalcevic (1980) presents a detailed discussion of different
types of flares, flare design and operating considerations, and a method
for estimating capital and operating costs for  flares.3  Elevated flares
are most common in the polymers and resins industry.  The basic elements
of an elevated flare system are shown in  Figures 3-1 and 3-2.  Process
offgases are sent to the flare through the collection header.  The
offgases entering the header can vary widely in volumetric flowrate,
moisture content, VOC concentration, and  heat value.  The knock-out drum
removes water or hydrocarbon droplets that could create problems  in the
flare combustion zone.  Offgases are usually passed through a water seal
before going to the flare.  This prevents a possible flame flashback,
which can be caused when the offgas flow  to the flare is too low  and  the
flame front moves down  into the stack.
      Purge gas  (Ng, COg, or natural gas)  also helps to prevent flashback
in the flare stack caused by low offgas flow.   The total volumetric flow
to the flame must be carefully controlled to prevent low flow flashback
problems and to avoid a detached flame  (a space between the stack and
flame with incomplete combustion) caused  by an  excessively high flowrate.
A gas barrier or a stack seal is sometimes used just below the flare
head  to  impede  the flow of air into the flare gas  network.

                                 3-3
-------
Emission
 Sowce •
  Gas
               Gas Collection Header
                and Transfer Line
1
                           Disentrainment
                                   Drum
                                               Steam
                                             Nozzles


                                               Flare
                                               Head
                                                                       Gas
                                                                    Barrier
                                                                                                         Pilot
                                                                                                         Burners
                                                                     Flare
                                                                     Stack
                                       Purge	,   I
                                        Gas
                                      Water
                                      Seal
                                                 Drain
                      Figure  3-1.   Steam Assisted  Elevated  Flare System
                                                                                                               Steam
                                                                                                               Line
4	Ignition
        Device
        Air Line
        Gas Line
                                                      3-4
-------
                                       PILOT AND
                                       MIXER
    PILOT
    ASSEMBLY

STEAM
HEADER,
                       INTERNAL
                       STEAM
                       INJECTOR
                       TUBES
STEAM
DISTRIBUTION
RING
      TIP SHELL
         PUN
                     CONTINUOUS
                     MUFFLER
                                         CENTER STEAM
                                         JET
                                     ELEVATION
  Figure  3-2.   Steam  Injection  Flare Tip
                     3-5
-------
     The VOC stream enters at the base of the flame where it is heated
by already burning fuel and pilot burners at the flare tip.  Fuel flows
Into the combustion zone where the exterior of the microscopic gas
                                                    i   !   '              I
pockets is oxidized.  The rate of reaction is limited by the mixing of
the fuel and oxygen from the air.  If the gas pocket has sufficient
oxygen and residence time in the flame zone it can be completely burned.
A diffusion flame receives its combustion oxygen by diffusion of air
into the flame from the surrounding atmosphere.  The high volume of fuel
flow in a flare requires more combustion air at a faster rate than
simple gas diffusion can supply, so flare designers add steam injection
nozzles to increase gas turbulence in the flame boundary zones, drawing
1n more combustion air and improving combustion efficiency.  The steam
injection promotes smokeless flare operation by minimizing the cracking
reactions that form carbon.  Significant disadvantages of steam usage
are the increased noise and cost.  The steam requirement depends on the
composition of the gas flared, the steam velocity from the injection
nozzle, and the tip diameter.  Although some gases can be flared smokelessly
without any steam, typically 0.15 to 0.5 kg of steam per kg of hydrocarbon
in the flare gas is required.
     Steam injection is usually  controlled manually with the operator
observing the flare (either directly or on a television monitor) and
adding steam as required to maintain smokeless operation.  Several flare
manufacturers offer devices which sense flare flame characteristics and
adjust the steam flowrate automatically to maintain smokeless operation.
     Some elevated flares use forced air instead of steam  to provide the
combustion air and mixing required 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 annul us.  The principal  advantage
of air assisted flares is that  expensive steam is not required.  Air
assist  is rarely used  on large  flares because  air flow is  difficult to
control when the gas flow is  intermittent.  About 600 J/sec  (0.8 hp)  of
blower capacity is  required for  each 45  kg/hr  (100 Ib/hr)  of gas flared
(Klett and Galeski, 1976).4
                                  3-6
-------
     Ground flares are usually enclosed and have multiple burner heads
that are staged to operate based on the quantity of gas released to the
flare.  The energy of the gas itself (because of the high nozzle pressure
drop) is usually adequate to provide the mixing necessary for smokeless
operation and air or steam assist is not required.  The fence or other
enclosure reduces noise and light from the flare and provides some wind
protection.
     Ground flares are less numerous and have less capacity than elevated
flares.  Typically they are used to burn gas "continuously" while steam
assisted elevated flares are used to dispose of large amounts of gas
released in emergencies (Payne, 1982).5
     3.1.1.1  Flare VOC Destruction Efficiency.  The flammability limits
of the gases flared influence ignition stability and flame extinction
(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.  Outside the flame, so much air may be induced
that the flame is extinguished.  Fuels with wide limits of flammability
are therefore usually easier to burn (for instance, H2 and acetylene).
However, in spite of wide flammability limits, CO is difficult to burn
because it has a low heating value and slow combustion kinetics.
     The auto-ignition temperature of a fuel affects combustion because
gas mixtures must be at high enough temperature and at the proper mixture
strength to burn.  A gas with a low auto-ignition temperature will
ignite and burn more easily than a gas with a high auto-ignition temperature.
Hydrogen and acetylene have low auto-ignition temperatures while CO has
a high one.
     The heating value of the fuel also affects the flame stability,
emissions, and structure.  A lower heating value fuel produces a cooler
flame which does not favor combustion kinetics and also is more easily
extinguished.  The lower flame temperature will also reduce buoyant
forces, which reduces mixing (especially for large flares on the verge
of smoking).  For these reasons, VOC emissions from flares burning gases
with low heat content may be higher than those from flares which burn
high heat content gases.
                                 3-7
-------
     Some fuels, also, have chemical differences (slow combustion kinetics)
sufficient to affect the VOC emissions from flares.  For instance, CO is
difficult to ignite and burn, and so flares burning fuels with large
amounts of CO may have greater VOC emissions than flares burning pure
VOC.
     The density of the gas flared also affects the structure and stability
of the flame through the effect on buoyancy and mixing.  The velocity in
many flares is very low, and, therefore, most of the flame structure is
developed through buoyant forces on the burning gas.  Lighter gases thus
tend to burn better, all else being equal.  The density of the fuel also
affects the minimum purge gas required to prevent flashback and the
design of the burner tip.
     Poor mixing at the flare tip or poor flare maintenance can cause
smoking (particulate).  Fuels with high carbon-to-hydrogen ratios (greater
than 0.35) have a greater tendency to smoke and require better mixing if
they are to be burned smokelessly.
     The following review of flares and operating conditions summarizes
five studies of flare combustion efficiency.
     Palmer (1972) experimented with a 1/2-inch ID flare head, the tip
of which was located 4 feet from the ground.   Ethylene was flared at 15
to 76 m/sec (50 to 250 ft/sec) and 0.12-0.62 x 106 J/sec (0.4-2.1 x 106
Btu/hr) at the exit.  Helium was added to the ethylene as a tracer at 1
to 3 volume percent and the effect of steam injection was investigated
in some experiments.  Four sets of operating conditions were investigated;
destruction efficiency was measured as greater than 99.9 percent for
three sets and 97.8 percent for the fourth.  The author questioned the
validity of the 97.8 percent result due to possible sampling and analytical
errors.  He recommended further sampling and analytical techniques
development before conducting further flare evaluations.
     Siege! (1980) made the first comprehensive study  of a commercial
flare system.   He studied burning  of refinery gas  on  a commercial flare
head manufactured by Flaregas Company.  The flare gases used consisted
primarily of hydrogen  (45.4 to 69.3 percent by volume) and light paraffins
(methane to butane).  Traces of HS were also present  in some  runs.  The
                                  3-8
-------
flare was operated with from 130 to 2,900 kilograms of fuel/hr (287 to
6,393 Ib/hr), and the maximum heat release rate was approximately 68.9 x 106
J/sec (23-5 x 106 Btu/hr).  Combustion efficiency and local burnout was
determined for a total of 1,298 measurement points.  Combustion efficiency
was greater than 99 percent for 1,294 points and greater than 98 percent
for all  points except one, which had a 97 percent efficiency.  The
author attributed the 97 percent result to excessive steam addition.
     Lee and Whipple (1981) studied a bench-scale propane flare.8  The
flare head was 2 inches in diameter with one 13/16-inch center hole
surrounded by two rings of 16 1/8-inch holes, and two rings of 16 3/16-inch
holes.  This configuration had an open area of 57.1 percent.  The velocity
through the head was approximately 1 m/sec (3 ft/sec) and the heating
rate was 0.09 x 106 J/sec (0.3 x 105 Btu/hr).  The effects of steam and
crosswind were not investigated in this study.  Destruction efficiencies
were greater than 99 percent for three of four tests.  A 97.8 percent
result was obtained in the only test where the probe was located off the
center!ine of the flame.  The author did not believe that this probe
location provided a valid gas sample for analysis.
     Howes, et al. (1981) studied two commercial flare heads at John Zink's
flare test facility.9  The primary purpose of this test (which was
sponsored by the EPA) was to develop a flare testing procedure.  The
commercial flare heads were an LH air assisted head and an LRGO (Linear
Relief Gas Oxidizer) head manufactured by John Zink Company.  The LH
flare burned 1,045 kg/hr (2,300 Ib/hr) of commercial propane.  The exit
gas velocity based on the pipe diameter was 8.2 m/sec (27 ft/sec) and
the firing rate was 12.9 x 106 J/sec (44 x 106 Btu/hr).  The LRGO flare
consisted of three burner heads 1 meter (3 feet) apart.  The three-burner
combination fired 1,909 kg/hr (4,200 Ibs/hr) of natural gas.  This
corresponds to a firing rate of 24.5 x 105 J/sec (83.7 x 106 Btu/hr).
Steam was not used for either flare, but the LH flare head was in some
trials assisted by a forced draft fan.  In four of five tests, combustion
efficiency was determined to be greater than 99 percent when sampling
height was sufficient to ensure that the combustion process was complete.
One test resulted in combustion efficiency as low as 92.6 percent when
the flare was operated under smoking conditions.
                                 3-9
-------
     An excellent detailed review of the above four studies was done by
Payne, et al. in January 1982,l® and a fifth study  [McDaniel,  et al.
(1982)] determined the influence on flare performance of mixing, heat
content, and gas flow velocity.H  A summary of these studies is given
in Table 3-1.  Steam assisted and air assisted flares were tested at the
John Zink facility using the procedures developed by Howes.  The test
was sponsored by the Chemical Manufacturers Association (CMA) with the
cooperation and support of EPA.  All of the tests were with an 80 percent
propylene, 20 percent propane mixture diluted as required with nitrogen
to give different Btu/scf values.  This was the first work which determined
flare efficiencies at a variety of "nonideal" conditions where lower
efficiencies had been predicted.  All previous tests were of flares
which burned gases that were very easily combustible and did not tend to
soot.  This was also the first test that used the sampling and chemical
analysis methods developed for the EPA by Howes.
     The steam assisted flare was tested with exit flow velocities up to
19 m/sec (62.5 ft/sec), with heat contents of 11 to 81 x 106 J/scm  (294 to
2,183 Btu/scf) and with steam-to-gas (weight) ratios varying from zero
(no steam) to 6.86:1.  Flares without assist were tested down to 7.2 x 10^
J/scm (192 Btu/scf).  All of these tests, except for those with very
high steam-to-gas ratios, showed combustion efficiencies of over 98 percent.
Flares with high steam-to-gas ratios (about 10 times more steam than
required for smokeless operation) had lower efficiencies (69 to 82  percent)
when combusting 81 x 106 J/scm (2,183 Btu/scf) gas.
     The air assisted flare was tested with flow velocities up to 66 m/sec
(218 ft/sec) and with Btu contents of 3.1 to 81 x 106 J/scm (83 to  2,183 Btu/scf),
Tests at 10.5 x 10^ J/scm (282 Btu/scf) and above gave over 98 percent
efficiency.  Tests at 6.3 x 106 J/scm (168 Btu/scf) gave 55 percent
efficiency.
     After consideration of the results of these five tests, EPA
concluded that 98 percent combustion efficiency can be achieved by
steam assisted flares when these flares are operated with combustion
gas heat contents and exit flow velocities within ranges determined by
                                                                       i
the tests.  Under the tests conducted, steam flares were shown to
obtain 98 percent combustion efficiency combusting gases with heat
contents over 11.2 x 106 J/scm (300 Btu/scf) at velocities of
                                3-10
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less than 18.3 m/sec (60 ft/sec).  Steam flares are not normally operated
at the very high steam-to-gas ratios that resulted in low efficiency  in
some tests because steam is expensive and operators make every effort to
keep steam consumption low.  Flares with high steam rates are also noisy
and may be a neighborhood nuisance.  Nonassisted pipe flares were shown
to obtain 98 percent efficiency with heat contents over 200 Btu/scf at
velocities of less than 18.3 m/sec (60 ft/sec).  Air assisted flares
were shown to obtain 98 percent efficiency with heat contents over 11.2
x 106 0/scm and at velocities not exceeding that determined by the
following formula:
           v(ft/sec) « 28.75 + 0.867 HC
           where v   = maximum gas velocity in ft/sec, standard conditions,
                HC   = heat content of the combusted gas in Btu/scf.
     The EPA has a program underway to determine more exactly the efficiencies
of flares used  in the petroleum/SOCMI industries and a flare test facility
has been constructed.  The combustion efficiency of  four flares (3.8 to
30.5 cm d1a.) will be determined and the effect on efficiency of  flare
operating parameters, weather factors, and heat content will be established.
The efficiency  of larger  flares  will be  estimated  by scaling.  A  final
report of this  work  should be available  in the spring of 1984.
     3.1.1.2  Applicability  of Flares.   A typical  polymer  plant produces
several hundred million pounds of  product per year.  Because of this
huge throughput, the VOC  emissions that  result from  frequent process
upsets are  also large.  Flares are used  mainly to  minimize the  safety
risk caused by  emergency  blowdowns from  high  pressure  processes where
large  volumes of  gases  with  variable composition  must  be released from
the plant  almost  instantaneously.   Flares are  ideal  for  this  service  and
their  reliability,  as measured by  absence of  explosions  and plant fires*
has been demonstrated repeatedly.   Flares also  effectively eliminate  the
hazard of  process  streams which, during  startup or shutdown, would
otherwise  vent  to the atmosphere and  could  also create an  explosion  or
toxic  hazard.   Finally,  flares are also  used to burn co-products  or
by-products of  a  process  that has  too  little value to reclaim,  and thus
would  otherwise be a continuous  VOC emission during normal operation  of
the unit.   This practice, which  was the standard  practice  for low pressure
processes  such  as the liquid phase polypropylene  and polyethylene processes,
                              •   3-12
-------
has become less common during the past decade as the value of VOC stream
components has dramatically increased.
3.1.2  Thermal Incinerators
     The design and operation of thermal incinerators are influenced by
operating temperature, residence time, desired VOC destruction efficiency,
offgas characteristics, and combustion air.  Operating temperatures may
typically be between 650°C (1,200°F) and 980°C (1,800°F) with a residence
time of 0.3 to 1.0 second.12  The temperature theoretically required to
achieve complete oxidation depends on the nature of the chemical involved
and can be determined from kinetic rate studies.13  The design of the
combustion chamber should maximize the mixing of the VOC stream, combustion
air, and hot combustion products from the burner.  This helps ensure
that the VOC contacts sufficient oxygen while at combustion temperature,  •
for maximum combustion efficiency.
     The heating value and water content of the waste gas feed and the
excess combustion air delivered to the incinerator also affect incinerator
design and operation.  Heating value is a measure of the heat produced
by the combustion of the VOC in the waste gas.  Gases with a heating
value less than 1,860 kJ/scm (50 Btu/scf) will not burn and require
auxiliary fuel to maintain combustion.  Auxiliary fuel requirements can
be reduced and sometimes even eliminated by transferring heat from the
exhaust gas to the inlet gas.  Offgases with a heating value between
1,860 kJ/scm  (50 Btu/scf) and 3,720 kJ/scm (100 Btu/scf) can support
combustion but require some auxiliary fuel to ensure flame stability,
i.e., avoid a flameout.  Theoretically, offgases with a heating value
above 3,720 kJ/scm (100 Btu/scf) possess enough heat content to not
require auxiliary fuel (although practical experience has shown that
5,580 kJ/scm  (150 Btu/scf) and above may be necessary)14 and these
offgases may be used as a fuel gas or boiler feed gas.I5  A thermal
incinerator handling offgas streams with varying heating values and
moisture content requires periodic adjustment to maintain the proper
chamber temperatures and operating efficiency.  Increases in heat
content reduce auxiliary fuel requirements, whereas  increases in water
content can substantially increase fuel requirements.
      Incinerators are always operated with excess air to ensure a  sufficient
supply of oxygen.  The amount of excess air used varies with the fuel
                                 3-13
-------
and burner type but should be kept as low as possible.   Using too much
excess air wastes fuel  because this air must be raised  to the combustion
temperature but does not contribute any heat by participating in the
oxidation reaction.  Large amounts of excess air also increase the flue
gas volume and may cause an operator to invest in a larger system than
required.
     A thermal incinerator usually contains a refractory-lined chamber
(which may vary in cross-sectional size along its length) containing a
burner at one end.  Because of the risk to the refractory, incinerators
are neither brought quickly up to nor cooled down quickly from operating
temperatures.  They require a fairly constant fuel input to maintain
combustion temperature.  A diagram of a thermal incinerator using discrete
burners is shown in Figure 3-3.  (Numbers in parentheses following the
mention of equipment parts or streams denote the numbered items on the
referenced figures.)  Discrete dual fuel burners (1) and inlets for the
Offgas (2) and combustion air (3) are arranged in a premixing chamber
(4) to thoroughly mix the hot products from the burners v.'ith the off gas
air streams.  The mixture of hot reacting gases then passes into the
main combustion chamber (5).  This section is sized to allow the mixture
enough time at the elevated temperature for the oxidation reaction to be
completed (residence times of 0.3 to 1 second are common).  Energy can
then be recovered from the hot flue gases with the installation of a
heat recovery section  (6).  Preheating of combustion air or the process
waste offgas fed to the incinerator by the incinerator exhaust gases
will reduce auxiliary  fuel usage.  In  some instances, the incinerator
exhaust gas may be used in a waste heat boiler to generate steam.
Insurance regulations  require that if  the process waste  offgas is preheated,
the VOC concentration  must be maintained below 25 percent of the lower
explosive limit  (LEL)  to minimize  explosive  hazards.^
     Thermal incinerators designed specifically for  VOC  incineration
with natural  gas as the auxiliary  fuel may use a  grid-type (distributed)
gas burner similar to  that shown  in Figure 3-4.  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  for mixing  the gases  entering the
chamber  (3).  This arrangement ensures burning of all  vapors  using less
fuel and a shorter burning length  in the duct than conventional  forward
                                 3-14
-------
  Waste Gas
  Auxiliary
Fuel Burner
  (discrete)
  Waste Gas
                                                                            Stack
                    Mixing
                   Section
Combustion
  Section
Optional
  Heat
Recovery
            Figure 3-3.   Discrete Burner Thermal  Incinerator
                         Burner Plate-, .  Flume Jets
                                                7
                                   Stack
                                                                  Optional
                                                                    Heat
                                                                  Recovery
                         (natural gas)
                         Auxiliary Fuel

         Figure 3-4.   Distributed Burner  Thermal  Incinerator
                                       3-15
-------
flame burners.  Overall, this system makes possible a shorter reaction
chamber while maintaining high efficiency.1?
     Thermal incinerators used to burn halogenated VOC's often use
additional equipment to remove the corrosive combustion products.   The
flue gases are quenched to lower their temperature arid routed through
absorption equipment such as spray towers or liquid jet scrubbers  to
remove the corrosive gases from the exhaust.18
     Packaged, single unit thermal incinerators are available in many
sizes to control streams with flowrates from a few hundred scfm' up to
about 50,000 scfm.  A typical thermal incinerator built to handle  a VOC
waste stream of 850 scm/min (30,000 scfm) at a temperature of 870°C
(1,600°F) with 0.75 second residence time would probably be a refractory-
lined cylinder.  With the typical ratio of flue gas to waste gas of
about 2.2, the chamber volume necessary to provide for 0.75 second
residence time at 870°C (1,600°F) would be about 100 m3 (3,500 ft3).  If
the ratio of the chamber length to the diameter is 2, and if a 30.5 cm
(1 ft) wall thickness is allowed, the thermal incinerator would measure
8.3 m (27 ft) long by 4.6 m (15 ft) wide, exclusive of heat exchangers
and exhaust equipment.
     3.1.2.1  Thermal Incinerator VOC Destruction Efficiency.  The
destruction efficiency of an incinerator can be affected by variations
in chamber temperature, residence time, inlet concentration, compound
type, and flow regime (mixing).  Of these, chamber temperature, residence
time, and flow regime are the most important.
     When the temperature exceeds 700°C (1,290°F), the oxidation reaction
rate is much  faster than the rate at which mixing can take place,  so VOC
                                                  -              •'         !
destruction becomes more dependent upon the  fluid mechanics within the
combustion  chamber.19   Variations in  inlet concentration also affect the
VOC destruction efficiency achievable; kinetics calculations describing
the combustion reaction mechanisms indicate  much slower reaction rates
at very low compound concentrations.  Therefore, at low VOC concen-
tration,  a  greater residence time is  required to achieve a high combustion
efficiency.
     Test results show  that  a VOC control efficiency of 98 percent can
be achieved consistently for many VOC compounds by well-designed units
and can be met under a  variety  of operating  conditions:^,21 combustion
                                 3-16
-------
chamber temperatures ranging from 700 to 1,300°C (1,300 to 2,370°F)  and
residence times of 0.5 to 1.5 seconds.  The test results covered the
following VOC compounds:  C-| to C$ alkanes and olefins, aromatics (benzene,
toluene, and xylene)s oxygenated compounds (methyl  ethyl ketone and
isopropanol), chlorinated organics (vinyl  chloride), and nitrogen-containing
species (acrylonitrile and ethylamines).  Although a combustion chamber
temperature of 870°C (1600°F) and a residence time of 0.75 seconds was
chosen for the cost analysis, the test results show that 98 percent
destruction efficiency is sometimes available at temperatures of 700°C
(1300°F) and residence times of 0.5 to 1.5 seconds.20
     Based on the studies of thermal  incinerator efficiency, auxiliary
fuel use, and costs, EPA has concluded that 98 percent VOC destruction,
or a 20 parts per million by volume (ppmv) compound exit concentration
(whichever is less stringent), is the highest reasonable control level
achievable by all new incinerators considering current technology.22
     3.1.2.2  App.1.1 cabi 1 _1 ty of Therma 1 Incinerators.  Thermal incinerators
can be used to control a wide variety of continuous waste gas streams (one
has been observed in a polypropylene plant23).  They can be used to
destroy VOC in streams with any concentration and type of VOC.  Although
they accommodate minor fluctuations in flow, incinerators are not well
suited to streams with intermittent flow because of the large -auxiliary
fuel requirements during periods when there is no fuel contribution from
the waste gas, yet the chamber temperature must be maintained to protect
the incinerator lining.
     For extremely dilute streams, a catalytic incinerator might be a
favorable choice over a thermal incinerator if supplemental fuel requirements
are of principal concern.   However, most waste gas streams in this
industry contain enough heating value to support a flame by itself on a
properly designed flame burner.  Such streams can be considered for use
as fuel gas or boiler feed  gas, from which the recovery of energy may
more than compensate for a  thermal incinerator's capital costs.
3.1.3  Catalytic Incinerators
     The control principles and equipment used in catalytic incineration
are similar to those employed in conventional thermal  incineration.  The
VOC-containing waste gas stream is heated to an appropriate reaction
temperature and then oxidation is carried out at active sites on the
                                3-17
-------
surface of a solid catalyst.  The catalyst increases  the rate of  oxidation,
allowing the reaction to occur at a lower temperature than  in thermal
Incineration.  This technique may offer advantages over thermal  incineration
in auxiliary fuel savings where low VOC content makes large fuel  usage
necessary.  Catalytic incinerators also may produce less NOX because of  ,
lower combustion temperatures and smaller excess air  requirements.
     Combustion catalysts are made by depositing platinum or platinum
alloys, copper oxide, chromium, or cobalt on an inert substrate,  which
is suitably shaped to fit the mechanical design of the incinerator.   The
                                                                        !
operating temperature of the catalyst is usually from 315°C (600°F)  to   .
650°C (1,200°F).  Combustion may not occur below 315°C (600°F) and
temperatures higher than 650°C (1,200°F) may shorten  the catalyst life
or even evaporate catalyst  from the support substrate.24  Accumulation
of particulate matter, condensed VOC's, or polymerized hydrocarbons  on
the catalyst can block the  active sites and reduce its effectiveness.
Catalysts can also be contaminated and deactivated by compounds containing
sulphur, bismuth, phosphorous, arsenic, antimony, mercury, lead, zinc,
tin, or halogens.  If the catalyst is so  "poisoned,"1 VOC's will pass
through unreacted or only partially oxidized.  Catalytic incinerators
can operate efficiently treating offgas streams with VOC concentrations
below the lower  explosive limit.  This  is  a distinct advantage over
thermal incinerators which  would in this  situation require auxiliary    ;
fuel.
     A  schematic of  a catalytic  incinerator unit  is  shown  in  Figure 3-5.
During  operation, the waste gases  (1) first enter the  mixing  chamber
 (also  called the preheat zone)  (3) where  they  are  heated by  contact with
the  hot combustion  products of a  burner (2).   The  mixing chamber temperature
may  vary  as a  function  of the composition and  type of  contaminants  to be
oxidized, but  will  generally operate  in the range  of 343°C  (650°F)  to
 593°C  (1,100°F).25   The heated mixture  then passes through  the catalyst
 bed  (4) where  oxygen and VOC's diffuse  to the  catalyst and  are adsorbed
 on its  surface.   The oxidation  reaction takes  place  at these "active
 sites."  Reaction products  desorb from the active  sites and  diffuse back
 into the  waste gas.   As with the exhaust gases from  thermal  incinerators,
 the products of combustion  leaving the bed may be used in  a waste heat
 recovery  device (5)  before  being exhausted to  the atmosphere.
                                 3-18
-------
3-19
-------
     3.1.3.1 Catalytic Incinerator VOC Destruction  Efficiency.   The
destruction efficiency of catalytic incinerators  is a function  of many
variables, including type of catalyst, its surface  area,  volume, and
pore size distribution, gas composition,  uniformity of flow through the
catalyst bed, oxygen concentration, and temperature in the unit.26»27
     The efficiency of a catalytic incinerator will deteriorate over
time, necessitating periodic replacement of the catalyst.  The  replace-
ment time varies widely, depending on the service of the  unit,  from less
than 1 year up to 10 years,12 with an average life between 3 and 5 years.28
     A 1980 study by Engelhard Industries for the EPA involved  testing
of both pilot and full-scale catalytic incineration systems. The  full-scale
unit installed on a formaldehyde plant achieved control efficiencies
ranging from 97.9 to 98.5 percent.  These efficiencies represent overall
control levels for carbon monoxide, methanol, dimethyl ether, and  formaldehyde.
Measurements indicated the ability of the system to control at  this
                                                                        j
level consistently over a 1-year period.  No trend in the data  points
gave indication of a maximum catalyst life.29
     3.1.3.2  Applicability of Catalytic  Incinerators.  A catalytic
incinerator  is best applied to a continuous stream that is (1)  low in
VOC  (higher  VOC concentrations lead to higher catalyst temperatures,
which can  seriously damage the catalyst activity and possibly create
fire hazards) and  (2)  free from solid particles and catalyst "poisons."
A catalytic  incinerator in many situations may be  favored over a thermal
incinerator  because it can destroy the VOC at a lower temperature and,
therefore,  use less fuel.  However,  since most of  the  streams  involved
in the  polymers and resins industry  are high enough in heating value to
self-combust without  using auxiliary  fuel, virtually  no advantage is
achieved  by  using  a catalytic unit and their applicability in this
industry  is  very limited.
3.1.4   Industrial  Boilers
     Fireboxes of  boilers  and fired  heaters  can  be used,  under proper
conditions,  to incinerate  waste streams that contain  VOC's.  Combustible
contaminants,  including  smoke, organic  vapors, and gases  can be converted
essentially to carbon dioxide and  water  in boiler  fireboxes.   As the
primary purpose  of the boiler is  to  generate steam,  all  aspects of
operation must be  thoroughly evaluated  before  this method of air pollution
                               3-20
-------
control can be used.  Any breakdown in the boiler can result in expensive
process downtime.  Consequently, the risk of shutdown should be kept
small and only streams that do not threaten boiler performance should be
introduced.
     For the satisfactory use of boilers as a control device, there are
several prerequisites.  Generally, the burner must be modified, the
boiler must operate continuously and concurrently with the pollution
source, the contaminants must be completely combustible, and the products
of combustion must not corrode the materials used to construct the
boiler.  Corrosive VOC compounds can be combusted in a boiler, but
special attention must be given to operate above the dew point of the
flue gases.  If these gases are allowed to condense, severe corrosion
problems will occur.  Further, the volumetric flowrate of low VOC concen-
tration emission streams must be taken into consideration because they
can reduce thermal efficiencies in the same way as excess combustion air
does.  The pressure drop caused by additional products of combustion
should not exceed the draft provided by boiler auxiliaries.  Boiler
life, efficiency, and capacity can be affected by the presence of con-
taminants in the VOC emission streams.  Halogens, for example, would be
devastating to the life of boiler tubes.  Finally, a personnel safety
hazard may occur if coal-fired boilers that are not pulverized coal-
fired are used to destroy organic waste.  Any interruption in the air
supply to these types of boilers would release into the boiler house
combustion vapors and any hazardous or toxic substances that may have
been injected.30  Great care, therefore, must be exercised in selecting
this mode of pollution control.
     The large majority of industrial boilers are of water tube design.
Water, circulated through the tubes, absorbs the heat of combustion.
Drums store the superheated water from which steam is directed to external
heat exchangers for use as process steam.  Boilers typically operate at
combustion chamber temperatures above 1,650°C (3,000°F) with a residence
time of about 1 second.31
     Both forced and natural  draft burners, designed to thoroughly mix
the incoming fuel and combustion air, may be used.  After ignition, the
mixture of hot reacting gases passes through the furnace section that is
sized to allow the oxidation  reaction to reach completion and to minimize
                                  3-21
-------
abrasion on the banks of the water tubes.  Energy transfer from the hot
flue gases to form steam can attain greater than 85 percent efficiency.
Additional energy can be recovered from the hot exhaust gases by instal-
lation of a gas-gas heat exchanger to preheat combustion air.
     Boilers designed specifically for use as a VOC control device
typically use discrete or vortex burners, depending on the heating value
of the vent stream.  For vent streams with heating values between
1,100 kJ/scm (300 Btu/scf) and 1,850 kJ/scm (500 Btu/scf), a discrete
burner would be best suited.31  Streams with lower heating values would
probably  require vortex burners to ensure the desired VOC destruction.
     3.1.4.1 Industrial Boiler VOC Destruction Efficiency.  VOC destruction
efficiency achievable by boilers depends on the same factors that affect
any combustion technique.  Since boiler furnaces typically operate at
higher peak temperatures and with longer combustion residence times than
thermal incinerators, the VOC destruction efficiency usually would be
expected  to match or exceed the 98 percent efficiency demonstrated in
incinerators.
     3.1.4.2  Applicability of Industrial Boilers.  Use of a boiler for
VOC emission control in the polymers and resins industry  is  uncommon.
Despite the potential problems, boilers  are being  usied  in at least two
polypropylene plants32  and  a high-density polyethylene  plant.33   The
polypropylene plants supplement boiler fuel with waste  gas that  otherwise
would  be  flared.  The high  density polyethylene  plant  sends  the  dehydrator
regeneration gas  (a  mixture of natural gas and nitrogen)  and a degassing
stream from the recycle diluent  step  (mostly  ethylene)  to steam-generating
boilers  as a fuel.
      A boiler would  be  used as  a  control device  only  if the  process
generated its  own  steam or  the  fuel  value  of  the waste  gas was  sufficient
to make the  process  a  net exporter of  steam.   Whenever either  condition
exists,  installation of a boiler  is  an  excellent control  measure that
provides greater than  98 percent  VOC  destruction and  very efficient
 recovery of  the heat of combustion of  the  waste gas,.
 3.2  CONTROL BY RECOVERY TECHNIQUES
      The three major recovery  devices  are  condenser;?,  adsorbers, and
absorbers.   These devices permit  many  organic materials to be  recovered
                                 3-22
-------
and, in some cases, reused in the process.  Condensers are widely used
for recovering organics from both continuous and intermittent rich
by-product streams in polystyrene manufacturing processes.  The VOC is
mainly styrene which is easily condensed because of its relatively high
condensation temperature.  The ease of styrene recovery and the ability
of a condenser to handle an intermittent stream makes it a desirable
control technology for all process VOC emissions in the polystyrene
industry.  Condensers may also be used in series with other air pollu-
tion control systems.  A condenser located upstream of an incinerator,
adsorber, or absorber will reduce the VOC load entering the downstream
control device.  The downstream device will abate most of the VOC that
passes through the condenser.
     Adsorbers are used on gas streams which contain relatively low VOC
concentrations.  Concentrations are usually well below the lower explosive
limit in order to guard against overheating of the adsorbent bed.
Adsorbers are often neither suitable nor the most efficient means of
control for the higher VOC concentration streams characteristic of the
polymers and resins industry.
     Absorbers, which use low volatility liquids as absorbents, are
another control option.  Their use is generally limited to applications
in which the spent absorbent can be used directly in a process, since
desorption of the VOC from the absorbent is often prohibitively expensive.
     Recovery techniques either condense the organic or contact the
VOC-containing gas stream with an appropriate liquid or solid.  Gases
containing only one or two organic gases are easier to process by recovery
techniques than multi-component mixtures.  The presence of inert or
immiscible components in the waste gas mixture complicates recovery
techniques.
3.2.1  Condensers
     Condensation devices transfer thermal energy from a hot vapor to a
cooling medium, causing the vapor to condense.  Condenser design thus
typically requires knowledge of both heat and mass transfer processes.
Heat may be transferred by any combination of three modes:  conduction,
convection, or radiation.
     The design of a condenser is significantly affected by the. number
and nature of components present in the vapor stream.  The entering
                                3-23
-------
gases may consist of a single condensable component or any number of
gaseous components which may or may not all  be condensable or miscible
with one another.  Example gas streams found in the polystyrene industry
may consist of a single condensable component (styrene);  a mixture of
condensable and noncondensable components (styrene and air); a mixture
of condensable, but immiscible, components (styrene and steam); or a
mixture of condensable, but immiscible, components v;ith a noncondensable
component (styrene, steam, and air).
     Condensers are designed and sized using the principles of thermodynamics.
At a fixed pressure, a pure component will condense isothermally at the
saturation or equilibrium temperature, yielding a pure liquid conden-
sate.  A vapor mixture, however, does not have a single condensate
temperature.  As the temperature drops, condensation progresses, and the
composition, temperature, enthalpy, and flowrate of both the remaining
vapor and the condensate will change.  These changes can be calculated
from thermodynamics data, if it is assumed that the vapor and liquid
condensate are in equilibrium.  Variations in composition and temperature
will affect most of the physical and transport properties which must be
used in condenser design calculations.  When these properties change,
the calculations governing the heat transfer process are adjusted to
accommodate these changes.
     In a two-component vapor stream with one noncondensable component,
condensation occurs when the partial pressure of the condensable component
is equal to the component's vapor pressure.  To separate the condensate
from the gas at fixed pressure, the temperature of the vapor mixture
must be reduced.  The liquid will begin to appear when the vapor pressure
of the condensable component becomes equal to its partial pressure, the
"dew point."  Condensation continues as the temperature is further
reduced.  The presence of a noncondensable component interferes with the
condensation process, because a layer of  noncondensabte on the condensate
acts as a heat transfer barrier.
     Two types of condensers are employed:  contact and surface.  Contact,
or direct, condensers cause the hot gas to mingle intimately with the
cooling medium.  Contact condensers usually operate by spraying a cool
liquid directly into the gas stream.  Contact condensers also may behave
as scrubbers since they sometimes collect noncondensable vapors which
                                 3-24
-------
are immiscible with the coolant.  The direct contact between the vapor
and the coolant limits the application of contact condensers since the
spent coolant can present a secondary emission source or a wastewater
treatment problem,34 unless it is economically feasible to separate the
two in a subsequent process.
     Surface, or indirect, condensers are usually common shell-and-tube
heat exchangers.  The coolant usually flows through the tubes and the
vapor condenses on the outside of the tubes.  In some cases, however, it
may be preferable to condense the vapor inside the tubes.  The condensate
forms a film on the cool tube and drains to storage.3^  The shell-and-tube
condenser is the optimum configuration from the standpoint of mechanical
integrity, range of allowable design pressures and temperatures, and
versatility in type of service.  Shell-and-tube condensers may be designed
to safely handle pressures ranging from full vacuum to approximately
41.5 MPa (6,000 psig), and for temperatures in the cryogenic range up to
approximately 1,100°C (2,000°F).36  Surface condensers usually require
more auxiliary equipment for operation (such as a cooling tower or a
refrigeration system) but offer the advantage of recovering valuable VOC
without contaminating the coolant, thereby minimizing waste disposal
problems.  The successively more volatile material returned from the
condenser to the distillation column is termed "reflux," or overhead
product.  The heavier compounds removed at the bottom are often called
column "bottoms."37
     The major pieces of equipment used in a typical refrigerated surface
condenser system are shown in Figure 3-6.3^  Refrigeration is often
required to reduce the gas phase temperature sufficiently to achieve low
outlet VOC concentrations.  This type of system includes dehumidification
equipment (1), a shell-and-tube heat exchanger (2), a refrigeration
unit (3), recovery tank  (4), and operating pumps  (5).  Heat transfer
within a shell-and-tube condenser occurs through several material layers,
including the condensate film, combined dirt and scale, the tube wall,
and the coolant film.  The choice of coolant used depends on the saturation
temperature of the VOC stream.  Chilled water can be used to cool down
to 4°C (40°F), brines to -34°C  (-30°F), and chlorofluorocarbons below
-34°C  (-30°F).39  Temperatures as low as -62°C (-80°F) may be necessary
to condense some VOC streams.34
                                3-25
-------
ii
                                                                 0)
                                                                 4->
                                                                 to
                                                                 (O
                                                                 in
                                                                 s:
                                                                 cu
                                                                 T3

                                                                 o
                                                                 o
                                                                 to

                                                                 oo
                                                                 O)
                          8
                                            UI
                                            Q:
                 3-26
-------
     3.2.1.1  Condenser Control  Efficiency.  VOC removal  efficiency of
a condenser is dependent upon the composition of the stream.   Single
component streams with a relatively high boiling point will  easily
condense, resulting in essentially 100 percent control efficiency.
Thus, very high efficiencies would be expected for condensers controlling
such streams in the polystyrene industry.  A less condensable component
in the stream, however, will reduce the control efficiency because of
the lower temperatures required for higher percentage removal.  Water-cooled
condensers sometimes cannot achieve a sufficiently low temperature to
ensure high control efficiency.  Better control, of course,  is possible
by use of a chilled coolant or even a refrigerated condenser at an
increased cost.  Outlet concentrations for low boiling organics may be
above 10,000 ppmv to 20,000 ppmv.40
          3.2.1.2  Applicability of Condensers.  Water-cooled condensers
are effective in reducing potential emissions of high boiling, easily
condensable organics, and find broad application in the polystyrene
manufacturing segment.  Surface condensers are used to recover styrene
from polystyrene vents.  Condensers cannot be used to condense low
boiling organics such as ethylene or propylene in streams containing
large quantities of inert gases such as nitrogen.  Refrigerated condensers
may be a viable option unless the stream contains water or heavy organics
which would freeze and foul the condenser.
3.2.2  Adsorbers
     Vapor-phase adsorption utilizes the ability of certain solids to
preferentially adsorb and thereby concentrate certain components from a
gaseous mixture onto their surfaces.  The gas phase (adsorbate) is
pumped through a packed bed of the solid phase (adsorbent) where selective
components are captured on its surface by physical adsorption.  The
organic molecules are retained at the surface of the adsorbent by means
of intermolecular or Van-der-Waals forces.  The adsorbed organics can be
readily removed and the adsorbent regenerated.
     The most common industrial vapor-phase adsorption systems use beds
of activated carbon.  Carbons made from a variety of natural  materials
(wood, coal, nut shells, etc.) are marketed for their special adsorbent
properties.  The multiple bed system maintains at least one bed online
while another is being regenerated.  Most systems direct the vapor

                              3-27  .
-------
_
           stream downward through a fixed carbon bed.  Granular carbon is usually
           favored because it is not easily entrained in the exhaust stream.
                Figure 3-7 is a schematic of a typical fixed bed, regenerative
           carbon adsorption system.  The process offgases are filtered and cooled (1)
           to minimize bed contamination and maximize adsorption efficiency.   The
           offgas 1s directed through the porous activated carbon bed (2) where
           adsorption of the organics progresses until the activated carbon bed is
           "saturated".  When the bed is completely saturated, the organic will
           "breakthrough" the bed with the exhaust gas and the inlet gases must
           then be routed to an alternate bed.  The saturated bed is then regenerated
                                                              .:•'•'   •    '   '     r | ,  '
           to remove the adsorbed material.
                Low-pressure steam  (3) is usually used to heat the carbon bed
           during the regeneration cycle, driving off the adsorbed organics,  which
           are usually recovered by condensing the vapors (4) and separating them
           from the steam condensate by decanting or distillation (5).  The adsorption/
                                                                   ;         .   ,     I
           regeneration cycle can be repeated numerous times, but eventually the
           carbon loses its adsorption activity and must be replaced.  The carbon
           can sometimes be reactivated by recharring.
                3.2.2.1  Adsorber Control Efficiency.  The efficiency of an adsorption
           unit depends on the properties of the carbon and the adsorbate, and on
           the conditions under which they contact.  Lower temperatures aid the
           adsorption process, while higher temperatures reduce the adsorbent's
           capacity.41  Removal efficiencies of 95 to 99 percent are achieved by
           well-designed and well-operated units.42
                3.2.2.2  Applicability of Adsorbers.  Adsorbers effectively control
           streams  with dilute concentrations of organics.  In fact, to prevent
           excessive temperatures within the bed due to the heat of adsorption,
           Inlet  concentrations of  organics are usually limited to  about 0.5 to
           1 percent.40  The maximum practical inlet concentration  is  about 1 percent,
           or 10,000 ppmv.43  Higher concentrations are frequently  handled by
           allowing some condensate to  remain from the  regeneration process to
           remove the heat  generated during adsorption.  Also,the  inlet stream can
           be diluted by use of a condenser or addition of  air or nitrogen upstream
           of the adsorber.  If the organic is reactive or oxygen is present in the
           vent stream, then additional precautions may be  necessary to safeguard
           the adsorption system.
                                            3-28
-------
VOC -Lateh
VentStraaffl
                                        121
 FILTERING
  AND
 COOLING
      (1)
OMn
   Low-pressure
   Steam
      (3)
                                        ADSORBER 1
                                        (ADSORBING)
                                       AOSOR^R 2
                                      (REGENERATING)
                                                     VENT TO
                                                   ATMOSPHERE
                                   Clond
                                                            Opan
                                                            (4)
                                                       I   CONDENSER   J
,
r
DECANTER
and/or
DISTILLING TOWER
Solvent
	 j» Water
                                                        (5)
           Figure  3-7.   Two Stage Regenerative Adsorption System
                                       3-29
-------
     Adsorbers can foul  and hence are not very suitable for streams
containing fine particles or polymerizeable monomers.   Both can  contaminate
the beds and result in poor performance, or even introduce safety  problems.
Because of their limitations in certain gas streams,  carbon adsorbers
are not ideally suited for most of the emission streams encountered  in
the polymers and resins industry.
3.2.3  Absorbers
       Absorption is a gas-liquid mass transfer operation in which a gas
mixture is contacted with a liquid (solvent) for the purpose of  preferentially
dissolving one or more components (solutes) of the gas.  Absorption  may
entail only the physical phenomenon of solution or may also involve
chemical reaction of the solute with constituents of the solvent.44
     For any given solvent, solute, and set of operating conditions,
there exists a theoretical equilibrium ratio of solute concentration in
the gas mixture to solute concentration in the solvent.  The driving
force for mass transfer in an operating absorption tower is related  to
the difference between the actual concentration ratio and this equilibrium
ratio.45  The solvents used are chosen for high solute (VOC) solubility
and include liquids such as water, mineral oil, nonvolatile hydrocarbon
oils, and aqueous solutions of oxidizing agents like sodium carbonate
and sodium hydroxide.46
     Devices based on absorption  principles include spray  towers, venturi
scrubbers, packed columns, and plate columns.  Spray towers and venturi
scrubbers are generally  restricted to  particulate removal  and control of
high-solubility gases.47  Most VOC control by  gas absorption is by
packed  or plate columns.  Packed  columns are  used mostly  for handling
corrosive materials,  liquids with foaming  or  plugging tendencies, or
where  excessive pressure drops would  result from the use  of  plate columns.
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).
Plate  columns are preferred for  large-scale operations, where internal
cooling is  desired, or  where low  liquid  flowrates would  inadequately  wet
the packing.48
     A schematic  of a packed tower  is  shown in Figure  3-8,  The gas is
introduced  at the bottom (1) and  rises through the packing material (2).
Solvent flows  by  gravity from  the top  of the  column  (3),  countercurrent
                                 3-30
-------
                                                                          CLEANED GAS OUT
                                                                       ^ To Final Control Device
ABSORBING
UQUIO IN
                                                                                       YOC LADEN
                                                                                       GAS IN
                                                    (4)
                                          ABSORBING LIQUID
                                           WITH VOC OUT
                                     To Disposal or VOC/Solvent Recovery
                      Figure  3-8.   Packed Tower for Gas Absorption
                                             3-31
-------
to the vapors, absorbing the solute from the gas phase and carrying the
dissolved solute out of the tower (4).  Cleaned gas exiting at the top
is ready for release or final treatment such as incineration.
     The major tower design parameters, column diameter and height,
pressure drop, and liquid flowrate, are based on the specific surface
area of the tower packing, the solubility and concentration of the
components, and the quantity of gases to be treated,,
     3.2.3.1  Absorber Control Efficiency.  The VOC removal efficiency
of an absorption device is very dependent on the characteristics of the
solvent and the design and operation  of the tower.  Generally, for a
given solvent and solute, an increase in absorber  size or  a decrease  in
the  operating temperature can  increase the VOC  control efficiency  of  the
system.
      Systems  that utilize  organic  liquids as  the solvent  usually  include
a separate item of  equipment to  strip the adsorbed gas so  that the
solvent can be  recycled  to the absorber.  The efficiency  of  the absorber
1s affected by  the  efficiency  of the  stripper.   For example,  a theoretical
 absorber calculated to achieve a removal  efficiency of 99.9  percent with
 once-through solvent usage (equivalent to 100 percent stripping efficiency),
 would achieve only  98.5 percent VOC removal  if the solvent were  recycled
 through a stripper which was 98 percent efficient.^9
      3.2.3.2  Applicability of Absorbers.  The selection  of absorption
 for VOC control depends on the availability of an appropriate solvent
 for the specific VOC.  Absorption is usually not considered when  the VOC
 concentration is below 200-300 ppmv.50
      The  use of absorbers is  generally limited to applications in which
 the stripped absorbent can  be reused directly  or  with minimum treatment.
 Absorption may not  be practical if the waste gas  stream contains  a
 mixture  of organics,  since  all will  likely not be highly  soluble  in  the
  same absorbent.  Absorbers  have found  limited  use as  a VOC  emission
  control  device in  the polymers  and resins  industry.
                                3-32
-------
3.3  REFERENCES FOR CHAPTER 3
1.   Lee, K.C., H.J. Jahnes, and D.C. Macauley.  Thermal Oxidation
     Kinetics of Selected Organic Compounds.  Journal of the Air
     Pollution Control Association.  29:749-751.  July 1979.  p. 750.

2.   Perry, R.H. and C.H. Chilton.  Chemical Engineers' Handbook, Fifth
     Edition.  McGraw-Hill Book Company.  1973.  p. 9-18.

3.   Kalcevic, V.  Emission Control Options for the Synthetic Organic
     Chemicals Manufacturing Industry, Control Device Evaluation, Flares
     and the Use of Emissions as Fuels.  U. S. Environmental Protection
     Agency.  Research Triangle Park, N.C.  Publication No.  EPA-450/3-80-026.
     December 1980.

4.   Klett, M.G. and J.B. Galeski.   Flare Systems Study.   Lockheed
     Missiles and Space Company.  NTIS Report PB-251664.   Publication
     No. 600/2-76-079.  March 1976.

5.   Payne, R., D. Joseph, J. Lee, C. McKinnon, and J.  Pohl.  Evaluation
     of the Efficiency of Industrial Flares Used to Destroy  Waste Gases.
     Phase I  Interim Report - Experimental  Design.  EPA Contract
     No. 68-02-3661.  Draft, January 1982.  p. 2-20.

6.   Palmer,  P.A.  A Tracer Technique for Determining  Efficiency of an
     Elevated  Flare.  E.I. duPont de Nemours and Company.   Wilmington,
     DE.   1972.

7.   Siege!,  K.D.   Degree of Conversion  of  Flare Gas  in Refinery High
     Flares.   University of Karlsruhe, The  Federal  Republic of  Germany*
     Ph.D. Dissertation.  February 1980.

8.   Lee,  K.C.  and  G.M. Whipple.  Waste  Gas Hydrocarbon Combustion  in  a
     Flare.   Union  Carbide  Corporation.   South  Charleston, W.V. (Presented
     at the  74th Annual Meeting of the Air  Pollution  Control  Association.
     Philadelphia,  PA.  June 21-26,  1981.)

9.   Howes,  J.E., T.E. Hill, R.N.  Smith,  G.R.  Ward,  and W.F. Herget.
     Development of Flare Emission Measurement Methodology.  EPA Contract
     No.  68-02-2682.   Draft,  1981.

10.  Reference 5,  p.  2-44 to 2-76.

11.  McDaniel, M.   Flare  Efficiency  Study,  Volume  I.   Engineering-Science.
     Austin,  Texas.   Prepared  for Chemical  Manufacturers  Association,
     Washington, D.C.   Draft  2,  January  1983.

12.  Kenson,  R.E.   A Guide  to  the Control  of  Volatile Organic Emissions.
     Systems Division,  MET-PRO Corporation.  Technical  Page 10T-1.
     Harleysville,  PA.   1981.
                                3-33
-------
13.  Reference 1, p. 749.

14.  Keller, M.  Comment on Control Techniques Guideline Document for
     Control of Volatile Organic Compounds Emissions from Manufacturing
     of High-Density Polyethylene, Polypropylene, and Polystyrene Resins.
     NAPCTAC Meeting.  June 1, 1981.  p. 6.

15.  Blackburn, J.W.  Organic Chemical Manufacturing, Volume 4:
     Combustion Control Devices, Report 1, Thermal Oxidation.  U. S.
     Environmental Protection Agency.  Research Triangle Park, N.C.
     Publication No. EPA-450/3-80-026.  December 1980.  p. 1-1.

16.  Basdekis, H.S.  Organic Chemical Manufacturing, Volume 4: Combustion
     Control Devices, Report 2, Thermal Oxidation Supplement (VOC
     Containing Halogens or Sulfur).  U. S. Environmental Protection
     Agency.  Research Triangle Park, N.C.  Publication No. EPA-450/3-80-026.
     December 1980.  p. 1-2 and 1-4.

17.  North American Manufacturing Company.  North American Combustion
     Handbook.  Cleveland, North American Mfg. Company.  1978.  p. 264.

18.  Reference 16, p. 1-1 and 1-2.

19.  Stern, A.C., ed.  Air Pollution, Third Edition, Volume IV, Engineering
     Control of Air Pollution.  New York, Academic Press.  1977.  p. 368.

20.  Mascone, D.C.  Thermal Incinerator Performance for NSPS.  U. S.
     Environmental Protection Agency.  Research Triangle Park, N.C.
     Memorandum to J.R. Farmer, Chemicals and Petroleum Branch.  June 11,
     1980.

21.  Mascone, D.C.  Thermal Incinerator Performance for NSPS, Addendum.
     U. S. Environmental Protection Agency.  Research Triangle Park,
     N.C.  Memorandum to J.R. Farmer, Chemicals and Petroleum Branch.
     July 22, 1980.

22.  Reference 20, p. 1.

23.  EEA, Incorporated.  Trip Report to ARCO Polymers, Inc.  EPA Contract
     No. 68-02-3061, Task 2.  1980.

24.  U. S. Environmental Protection Agency, Office of Air and Haste
     Management.  Control Techniques for Volatile Organic Emissions from
     Stationary Sources.  Research Triangle Park, N.C.  Publication
     No. EPA-450/2-78-022.  May 1978.  P. 32.

25.  U. S. Environmental Protection Agency, Office of Air and Water
     Programs.  Air Pollution Engineering Manual.  Research Triangle
     Park, N.C.  Publication  No. AP-40.  May 1973.  p. 180.

26.  Reference 25, p. 181.
                                3-34
-------
 27.  Reference 24, p. 34.

 28.  Key, J. A.  Organic Chemical Manufacturing, Volume 4:  Combustion
      Control Devices, Report 3, Catalytic Oxidation.  U. S. Environmental
      K? «S1/2no59n2^y'  Research Triangle Park, N.C.  Publication No.
      EPA-450/3-80-026.  December 1980.  p. 1 1-9.

 29.  Engelhard Industries Division, Engelhard Corporation.  Catalytic
      Incineration of Low Concentration Organic Vapors.  Prepared for
      M* s-  Environmental Protection Agency.  Washington, D.C.  Contract
      No. 68-02-3133.   January 1981.

 30.  Letter from Monsanto Company to J.R. Farmer, U. S.  Environmental
      Protection Agency.  April  19,  1982.   p. 17.

 31.  Memo from Senyk, D., EEA,  Inc., to Distillation file.

 32.  Shell  Chemical  Company,  Woodbury Plant.  Application for Permit to
      Construct, Install or Alter Control  Apparatus  or Equipment.   To New
             State Department  of Environmental  Protection.   March  16,
 33.   EEA,  Incorporated.   Trip  Report to Phillips  Chemical  Company.   EPA
      Contract  No.  68-02-3061,  Task  2.   August 8,  1980.

 34.   Erikson,  D.G.   Organic  Chemical  Manufacturing,  Volume 5:   Adsorption
      Condensation, and Absorption Devices,  Report 2,  Condensation.
      U.  S.  Environmental  Protection  Agency.   Research Triangle  Park,
      N.C.   Publication No. EPA-450/3-80-027.   December  1980.  p.  II-3.

 35.   Reference 24, p. 84.

 36.   Devore, A», G.J. Vago,  and 6.J.  Picozzi.   Heat  Exchangers:   Specifying
      and Selecting.  Chemical  Engineering.  87(20):133-148.  October  1980.
      p.  136.

 37.   Kern,  D.Q.  Process Heat  Transfer.  New  York, McGraw-Hill  Book
      Company.  1950.  p. 255.

 38.   Reference 34, p. 1 1-4.

 39.   Reference 34, p. IV-1.

 40.   Parmele, C.S.,  W.L. O'Connell, and H.S. Basdekis.  Vapor-Phase
      fSfSr?*1^" Cuts Po11ut1on» Recovers Solvents.  Chemical Engineerinq.
      86(28): 58-70.  December 1979.  p. 60.

41.  Basdekis,  H.S. and C.S.  Parmele.  Organic Chemical  Manufacturing,
     Volume 5:   Adsorption, Condensation, and Absorption Devices, Report  1
     Carbon Adsorption.   U. S.  Environmental Protection Agency.   Research
     Triangle Park, N.C.   Publication No. EPA-450/3-80-027.  December 1980
     p. II-l.
                               3-35
-------
42.  Reference 40, p. 69.
43.  Reference 40, p. 62.
44.  Reference 2, p. 14-2.
45.  Standifer, R.L.  Organic Chemical Manufacturing, Volume 5: .Adsorption,
     Condenstation, and Absorption Devices, Report 3, Gas Absorption.
     U. S. Environmental Protection Agency.  Research Triangle Park,
     N.C.  Publication No. EPA-450/3-80-027.  December 1980.  p. III-5.
46.  Reference 24, p. 76.
47.  Reference 45,  p. II-l.
48.  Reference 2, p. 14-10.
                                                            •            i
49.  Reference 45, p. III-6 and III-7.                                  '
50.  Reference 45, p. 1-1.                                              ',
                               3-36
-------
                  4.0  ENVIRONMENTAL ANALYSIS OF RACT

4.1  INTRODUCTION
     The environmental impact of the systems considered representative
of reasonably available control  technology (RACT) is essentially two-fold.
Primary impacts are those attributed directly to the control systems,
such as reduced levels of specific pollutants.  Secondary impacts are
indirect or induced in nature, such as aggravation of another pollutant
problem through utilization of a particular control system.  Both beneficial
and adverse environmental impacts related to each of the pollution
categories, air, water, and solid waste are assessed in the following
sections.  Also a discussion of the additonal amount and type of energy
required for control is included.
     The following emission reductions or limitations are considered
representative of RACT:
     (1)  For polypropylene plants using liquid phase processes: a 98
weight percent reduction or reduction to 20 ppm of continuous VOC
emissions from the polymerization reaction section (i.e., reactor
vents), the material recovery section (i.e., decanter vents, neutralizer
vents, by-product and diluent recovery operations vents), and the
product finishing section (i.e., dryer vents and extrusion and pelletizing
vents).
     (2)  For high-density polyethylene plants using liquid phase slurry
processes: a 98 weight percent reduction or reduction to 20 ppm of
continuous VOC emissions from the material recovery section (i.e.,
ethylene recycle treater vents) and the product finishing section (i.e.,
dryer vents and continuous mixer vents).
     (3)  For polystyrene plants using continuous processes: an emission
limit of 0.12 kg VOC/1,000 kg product from the material recovery section
(i.e., product devolatilizer system).
                                  4-1
-------
     These RACT recommendations were made on the basis of a single
control device controlling all vents of VOC emissions.  Many, if not
most, existing plants will have already some control that is representative
of RACT.  At these plants, additional emission control may be obtained
at the expense of installing an additional control device, which may be
controlling emissions from a single process section only.  The costs of
gaining additional emission reduction, therefore, depends upon the level
of control already employed at an existing plant.  For existing plants
that already have RACT level controls on most vents, a State may find
that control of the remaining vents may be very expensive in terms of
dollars per Mg of VOC emission reduction and the  installation of a
separate control device to reduce these emissions, while technically
feasible, may not be reasonable.
     Tables 4-1 through 4-3 present  emission rates  (and their equivalent
annual emissions) associated with various levels  of control cost
(expressed  as dollars per Mg  of VOC  reduction)  for  two potential control
situations.   (Appendix F  details the  derivation of  these  numbers.) The
first  situation is where  a  new single control device  is  needed  to
control the emissions from  one of the process sections in a  single
process line.  The second potential  control  situation is  where  a  new
single control device is  needed to  control the  emissions  from one of
the  process sections throughout the entire  plant  (i,,e.,  across  process
lines).   Using the emission rates  (or annual  emission levels) at  one  of
the  optional  cost-effectiveness cut-off levels  as a guideline,  States
may  exempt  plants with  uncontrolled emissions at  or below these emission
levels.   However,  it must be emphasized that the emission levels  in
Tables 4-1  through 4-3  were based  on a general  model  plant  and  used
approximate cost  equations.  Specific plants may have different stream
characteristics,  the  potential  to  combine streams,  or utilize  existing
control  devices that would make control of emissions at levels  below
those presented in Tables 4-1 through 4-3 reasonable.  Thus, the States
are encouraged to use a case-by-case approach for exempting any uncon-
trolled emission  stream from the  general  98 percent reduction (or to 20
 ppm reduction)  requirement.
      Combustion control devices such as flares, thermal  and catalytic
 incinerators, boilers,  and process heaters can  achieve 98 percent VOC
 destruction.   The following paragraphs discuss, as appropriate, the
                                 4-2
-------
Table  4-1.   UNCONTROLLED  EMISSION RATES VERSUS  COST EFFECTIVENESS
                     FOR  POLYPROPYLENE PLANTS BASED  ON  MODEL  PLANT
                           PARAMETERS,  BY  PROCESS SECTION
A. SINGLE PROCESS SECTION WITHIN
Process
Section
Polymerization
Reaction
Material Recovery
Product Finishing
Raw Materials
Preparation^
Control
Costs,
$/Mga
1,000
2,000
3,000
1,000
2,000
3,000
1,000
2,000
3,000
1,000
2,000
3,000
A SINGLE LINE (47 Gg capacity)
Uncontrolled Emission
kg VOC/Mg Product^
0.45
0.23
0.15
0.50
0.25
0.17
2.57
1.20
0.77
0.45
0.23
0.15
Rates
Mg/yrc
20.9
10.4
7
23.0
11.5
7.7
121
56
36
20.9
10.4
7
        B.   SINGLE PROCESS SECTION ACROSS LINES (141 Gg capacity)

                             Control         Uncontrolled Emission  Rates
Process
Section
Polymerization
Reacti on
Material Recovery
Product Finishing
Raw Materials
Preparation^
Costs,
$/Mga
1,000
2,000
3,000
1,000
2,000
3,000
1,000
2,000
3,000
1,000
2,000
3,000
kg VOC/Mg Product^
0.16
0.08
0.06
0.21
0.10
0.07
0.83
0.39
0.26
0.16
0.08
0.06
Mg/yrc
22.3
11.1
7.4
28.5
14.2
9.5
116
55
36
22.3
11.1
7.4
        Based on 98 percent reduction in the  uncontrolled emission  rate.
       b
        Emission rates have been rounded up to the nearest one-hundreth.

        Equivalent uncontrolled annual  emissions (Mg/yr) were calculated by
        multiplying the unrounded emission  rate (kg VOC/Mg product) times
        production capacity (i.e., 47 Gg or 141 Gg).
       d
        Typical emissions  from the raw materials preparation section are
        relatively small so that the cost of  emission reduction  is  considered
        unreasonable.   Therefore, control of  these emissions is  not part of
        the RACT recommendations.  However, emission levels may  vary from
        plant to plant. If raw materials preparation emissions  are at these
        levels or higher (as indicated  in the table), the State  may choose to
        require their  control.  Furthermore,  if raw materials preparation
        emissions can  be combined with  other  uncontrolled emissions (for
        example, from  the material  recovery section), then it is reasonable to
        control raw materials preparation emissions at even lower emission
        levels.


                                     4-3
-------
 Table 4-2.  UNCONTROLLED EMISSION RATES VERSUS COST EFFECTIVENESS
             FOR HIGH-DENSITY POLYETHYLENE PLANTS BASED ON MODEL
                     PLANT PARAMETERS, BY PROCESS SECTION
A. SINGLE PROCESS
Process
Section
Material Recovery
Product Finishing
B. SINGLE PROCESS
Process
Section
Material Recovery
Product Finishing
SECTION WITHIN
Control
Costs,
$/Mga
1,000
2,000
3,000
1,000
2,000
3,000
SECTION ACROSS
Control
Costs,
$/Mga
1,000
2,000
3,000
1,000
2,000
3,000
A SINGLE LINE (71.3 Gg capacity)
Uncontrolled Emission Rates
kg VOC/Mg Productb
0.30
0.15
0.10
0.87
0.41
0.27
LINES (214 Gg capacity)
Uncontrolled [Emission
kg VOC/Mg Productb
0.11
0.06
0.04
0.33
0.15
0.10
Mg/yrc
20.9
10.5
7
61.9
28.6
18.9
Rates
Mg/yrc
22.4
11.1
7.4
70
32.1
21.1
 Based  on  98 percent reduction  in  the  uncontrolled  emission  rate.

 Emission  rates have been  rounded  up to the  nearest one-hundreth.
•V                                                                     I
'Equivalent uncontrolled annual  emissions (Mg/yr) calculated by multiplying
 the unrounded emission rate (kg VOC/Mg product) times production
 capacity  (i.e., 71.3 Gg or 214 Gg).
                                  4-4
-------
  Table 4-3.   UNCONTROLLED EMISSION RATES VERSUS COST  EFFECTIVENESS
              FOR POLYSTYRENE PLANTS BASED ON MODEL PLANT PARAMETERS,
                              BY  PROCESS SECTION
A. SINGLE PROCESS
Process
Section
Material Recovery^
Material Recovery6
B. SINGLE PROCESS
Process
Section
Material Recovery^
Material Recovery6
SECTION WITHIN
Control
Costs,
$/M9a
1,000
2,000
3,000
1,000
2,000
3,000
SECTION ACROSS
Control
Costs,
$/Mga
1,000
2,000
3,000
1,000
2,000
3,000
A SINGLE LINE (36.75 Gg)
Uncontrolled Emission
kg VOC/Mg Product^
0.45
0.29
0.19
0.26
0.21
0.18
LINES (73.5 Gg)
Uncontrolled Emission
kg VOC/Mg Product^
0.20
0.17
0.16
0.19
0.17
0.15
Rates
Mg/yrc
16.4
10.7
7.1
9.3
7.5
6.7
Rates
Mg/yrc
14.6
12.5
11.5
13.7
11.9
11.1
 Based  on  the  emission  reduction  associated with  going  from  the
 uncontrolled  emission  rate  down  to  the  RACT  level  of 0.12 kg  VOC/Mg of
 product.
3
 Emission  rates  have  been  rounded up to  the nearest one-hundreth.

'Equivalent  uncontrolled annual emissions  (Mg/yr)  calculated by multiplying
 the  unrounded emission rate (kg  VOC/Mg  product)  times  production
 capacity  (i.e., 36.75  Gg  or 73.5 Gg).
i
 Styrene in  air.
a
 Styrene in  steam.
                                  4-5
-------
design and operating conditions that, based upon available data,  ensure
98 percent VOC destruction, the feasibility of emission testing,  and
the acceptability of existing units.
     A recent comprehensive flare emissions testing program conducted
jointly by EPA and the Chemical Manufacturers Association has demonstrated
that the following conditions ensure 98 weight percent VOC destruction:
smokeless operation (no visible emissions except for periods of 5 minutes
or less during a 2-hour period); the presence of a flame; a net heating
value of 300 Btu/scf or greater if the flare is steam-assisted or air-
assisted or of 200 Btu/scf or greater if the flare is non-assisted,  and
an exit velocity of 60 fps or less if steam-assisted or non-assisted  or
less than [8.706 + 0.7084 (Hr)] fps, where HT is the net heating  value,
                                           	•• ".     -it   I      i  ...     '   !
if the flare is air-assisted.  Operating conditions other than those  above
have not been investigated and there is no assurance that VOC destruction
efficiencies of 98 percent or greater would be achieved.  The high cost
makes it impractical to test a flare.  Therefore, a State may accept
new and existing flares as RACT provided the flares are operated  smoke-
lessly, with a flame, with minimum heat contents (Btu/scf) as outlined
above, and with maximum exit velocities as outlined above; except that
existing flares do not need to meet the maximum exit velocity guidelines
when major structural changes, such as flare tip replacement, are
required to meet the maximum exit velocity recommended for the particular
flare.
     For thermal incinerators, a control efficiency of 98 percent VOC
destruction or reduction to 20 ppm VOC, whichever  is less stringent,  is
considered to be achievable by all new incinerators considering available
technology, cost, and energy usage.  This determination is based on
considering incinerator operating conditions of 870°C  (1600°F), a
residence time of 0.75 seconds, and adjustment of  the  incinerator after
start-up.  As stated in Chapter 3, operating conditions other than
those noted above may still result in 98 percent emission reduction.
Thus, some existing incinerators designed and operated at lower combustion
temperatures and residence times may perform as well;  others may not.
An emission test of an incinerator is technically  and  economically
feasible.  Therefore, a State may require emission tests and, based on
                                  4-6
-------
the results and an analysis of cost effectiveness, require modifications
to improve efficiency or even replacement of an existing incinerator.
     Catalytic incinerators can be designed and operated to achieve
98 percent destruction; however, general parameters to assure performance
can not be specified because the required parameters vary with the
characteristics of the waste stream.  The performance of catalytic
incinerators can be tested at reasonable cost.  Therefore, as with
thermal incinerators, a State may require an emission test, modification,
or replacement of an existing catalytic incinerator.
     Boilers and process heaters used for VOC reduction must have the
VOC vent  stream introduced into the flame zone of the boiler or process
heater to assure high combustion efficiency.  Boilers and process heaters
with a design heat input capacity of 150 million Btu/hour or greater are
generally operated at temperatures and  residence times greater than 1095°C
(2,000°F) and 1 second, respectively.   Thus, the probability of very high
VOC reduction efficiency (i.e., 98 percent or a VOC reduction to 20 ppm)
is a near certainty.  The  achievement of 98 percent destruction efficiency
for boilers and process heaters with design heat input capacities less
than 150  million Btu/hour  is  not so certain.  Since performance tests
can be conducted at  reasonable  cost, however, a State may require testing
and, based on  a cost effectiveness analysis, subsequent modification or
even replacement to  improve  combustion  performance.
     Other control techniques,  such as  those utilizing condensation,
absorption and adsorption, can  be  designed and operated to achieve a
98 weight percent  VOC  reduction.   Any techniques  that achieve the same
degree of control  should be  considered  equivalent  to and  acceptable  as
RACT.
     The  0.12  kg  VOC/1,000 kg of  product  emission limit  for  polystyrene
continuous  processes is  based on the use  of condensers.   This level  is
 in  agreement  with the current emission  factors  reported  by the  Chemical
Manufacturers  Association.  The use of  process  changes or other control
techniques that  achieve  the same  degree of control  should be considered
 equivalent to  and acceptable as RACT.
     Although  many existing plants are  expected  to be  achieving RACT
 already,  these control  technique  guidelines  establish  uniform,  reasonably
 available state-of-the-art control for  existing  plants  in all  non-attainment
                                 4-7
-------
areas nationwide and provide information regarding VOC emissions and
their control in polypropylene liquid phase processes, high-density
polyethylene liquid phase slurry processes, and polystyrene continuous
processes.
     Control techniques guidelines and RACT are not established in this
document for other polymer processes, such as polypropylene gas phase
processes, polyethylene gas phase processes, and high-density polyethylene
liquid phase solution processes.  Emissions and control of these processes
were not analyzed because of the relatively small number of existing
plants.  EPA may subsequently analyze and establish control technique
                                                   , ,      , ,   ,  „,       , • j
guidelines for any or all of such other processes.  In the meantime, a
State may choose to conduct its own model plant or case-by-case analysis
and establish its own guidelines.
4.2  AIR POLLUTION
     The annual quantities of volatile organic compounds (VOC) from the
model plants before and after control by RACT are presented in Table 4-4.
The stream from each model plant represents a combination of continuous
emission streams from process vents excluding fugitives and raw material
and product  storage facilities.  The  range of expected reductions  in VOC
emissions, achieved as a result of implementation of  RACT for the  model
plants, is shown in Table 4-4.
     The VOC destroyed or recovered as a  result  of the application of
RACT consists mainly of ethylene, propylene, styrene, and certain
organic diluents.  These gases  are known  to  react in  the atmosphere
with oxides  of nitrogen to  form oxidants,  principally ozone.  Reduction
of  emissions of these gases will contribute to the attainment of the
national  ambient air quality standard (NAAQS)  for ozone.
     A flare is expected to be  the major  control  device used  as RACT  for
polypropylene  liquid phase  processes  and  polyethylene liquid  phase
slurry processes.   A properly designed  combustion device would  lead to
minimal formation  and subsequent emission of carbon monoxide.   The
amount of NOX  products  formed by  flaring  or  by  incineration  at  870°C
(1,600 °F)  is  negligible.   Thus, there  should  be minimal  generation of
secondary air  pollutants by combustion techniques.
                                 4-8
-------














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                                                                              	'""HBII'i ill	li'TV-il
     A condenser is expected to be the major control  device used as RACT
for polystyrene continuous processes.  There should be no generation of
secondary air pollutants by condensation.
4.3  WATER POLLUTION
     Combustion systems do not generate an effluent water stream.
(Boilers do generate an. effluent water stream during blowdown, but
combustion of VOC generates neither additional effluent nor changes in
effluent characteristics.)  The condensers for each polystyrene model
plant could require as much as 38,000 gallons of make-up water  (at a
cost of about $12 per year).  Most, of the condenser water losses, however,
would be expected to be by evaporation rather than by discharge.
4.4  SOLID WASTE DISPOSAL
     Generation of  solid wastes  is not an expected result of  control by
RACT in any model plant under consideration.  Relatively small  amounts
of used catalyst would  be  generated  if a catalytic incinerator  were  used
to separately  control some  of the low VOC streams.
4.5  ENERGY
     Tables 4-5, 4-6, and  4-7  present the additional  amount and type of
energy  required  after control within each model  plant by RACT.   The control
techniques analyzed for RACT are flares, thermal  incinerators,  and
catalytic  incinerators  for polypropylene and high-density  polyethylene
and  condensers for polystyrene.   These  control  techniques  require
steam,  natural  gas, and electricity.  Total  estimated energy consumption
 is presented  for each application in equivalent barrels of distillate
 oil  and total  cost.
      For a flare,  steam is generally used to ensure smokeless combustion.
 Natural gas is used for pilot flames to assure ignition and subsequent
 combustion of the waste gas.  The combined streams from the liquid-phase
 processes of both polypropylene and high-density polyethylene production
 have high enough heat contents that no  supplemental fuel is  required.
 In addition, for this same reason, no supplemental fuel is required for
 the control of emissions from individual process sections  for those
 process sections controlled by a  flare.  As the flow rate  of the flares
                                    4-10
-------






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would be largely, if not entirely, continuous, no natural  gas will  be
required as a purge.
     For an incinerator, no fuel is required for flame stability or
combustion since the streams encountered in these polymer industries are
rich enough to sustain stable self-combustion.  Electricity cost projected
is only for fan operation.  Instrumentation is assumed to consume a
negligible amount of electricity.  If natural gas were used as a supplemental
fuel, the possibility of fuel switching (gas to coal) is remote for an
incinerator.
     For a condenser, electricity is required to pump the cooling liquid.
Electricity may also be required to operate the refrigeration system of
a refrigerated condenser.  A condenser has no other energy requirements.
                                   4-14
-------
                  5.0  CONTROL COST ANALYSIS OF RACT
     This chapter presents assumptions, procedures, and results of the
analysis to estimate the costs of controlling volatile organic compounds
(VOC) emissions from the polymers and resins industry.  The results are
estimates of capital costs, annualized costs, and costs of emission
reductions, for a range of existing control  levels.  The following
sections present outlines of the bases for estimating control  costs for
flares, thermal incinerators, catalytic incinerators, and condensers
(more detailed procedures are given in Appendix E) and the results of
the cost analyses for each model plant.
5.1  BASES OF COST ANALYSES
     The cost analysis consists of two steps for each control  system:
designing a system that will reliably maintain the desired efficiency
and estimating capital and operating costs for such a system.   Designing
a control system for process VOC emissions requires an analysis of the
waste gas characteristics of the combined stream to each control device
specified for a model plant.  The various streams for each model plant
were assumed to have the same compositions assumed for the new source
performance standard cost analysis.1  The stream characteristics along
with mass and energy balances are the basis for determining the equipment
sizes, operating parameters, and operating requirements (e.g., fuel).
     Once these control system parameters have been determined, then the
capital and annual costs can be calculated.   The capital cost  estimates
for each control device and model plant combination include purchase and
installation of the control or monitoring devices and piping systems
necessary for proper control of continuous process VOC emissions from
each model plant.
     All process VOC control capital costs are converted to June 1980
dollars using the plant cost indices published in the Chemical Engineering
Economic Indicators.  The installed capital  costs for process  controls
represent the total investment, including indirect costs such  as engineering
and contractors' fees and overhead, required for purchase and  installation
of all equipment and materials for the control systems.  These are
                                  5-1
-------
battery-limit costs and do'not include any provisions for bringing
utilities, services, or roads to the site, or for any backup facilities,
land, research and development required, or for any process piping and
instrumentation interconnections that may be required within the process
generating the waste gas.  Since RACT will affect existing plants, the
control equipment installation factors include cost adjustments for
retrofit installations.  Typical cost adjustments for control equipment
installations given in the GARD Manual2 are presented in Table 5-1.  The
installation factors and retrofit cost adjustments assumed for the
various process control devices are presented in Table 5-2.  Actual
direct and indirect cost factors depend upon the plant specific conditions
and may vary with the size of the system.  The annualized costs consist
of the direct operating and maintenance costs, including labor, utilities,
fuel, and materials for the control system, and indirect costs for
overhead, taxes, insurance, administration, and the capital recovery
charges.  The utilities considered include natural gas and electricity.
The annualized cost factors that are used to analyze all of the process
VOC control systems are summarized in Table 5-3.
     The following sections outline the design and costing procedures
developed for.flares, thermal incinerators, catalytic incinerators,
and condensers.  Details of these procedures are given in Appendix E.
This section presents an overview of the procedures and their important
features.  The results of the cost analyses for the various control
device and model plant combinations are also presented.
5.1.1  Thermal  Incinerator Design and Cost Basis
     For costing purposes thermal incinerator designs were based on heat
and mass balances for combustion of the waste gas and any required
auxiliary fuel, considering requirements of total combustion air.
Associated piping, ducting, fans, and stacks were also costed.
     5.1.1.1  Thermal Incineration Design.  Designs of thermal incineration
systems for the various combinations of waste gas streams were developed
using a procedure based on heat and mass balances and the characteristics
of the waste gas in conjunction with some engineering design assumptions.
For the purpose of the cost analyses in this report, thermal incinerators
were designed to maintain a 0.75 second residence time at 870°C (1600°F).3
The design procedure is outlined in this section.
                                  5-2
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   Table 5-3.  ANNUALIZEO COST FACTORS FOR POLYMERS AND RESINS  CTG
                             (June 1980_ Dollars)

Direct Cost Factors
 Operating labor price: $18/hr (including overhead)9
 Operating labor requirements (incl uding. supervisory labor):
     « 1200 labor hours/yr for thermal  incinerator
            without heat recovery'3
     ~  620 labor hours/yr for flare0
     =   60 labor hours/yr for condenser^
     *  620 labor hours/yr for catalytic incinerator
                without heat recovery0
 Electricity price: $0.049/kwhe
 Natural gas price: $5.67/6J ($5.98/MMBtu)f
 Steam price: $13.62/Mg ($6.18/1000 1 b)9
 Water price: $0.079/m3 ($0.30/1000 gal )h
 Styrene recovery credit:  $0.788/kg ($0. 357/1 b)1'
Indirect Cost Factors
 Interest rates:
    10 percent  (in the absence of taxes)
 Equipment life, N:J
    15 years for flare
    10 years for thermal incinerator, catalytic incinerator,
       condenser, piping
Capital  recovery charge factor
        0.131 for flare
                               =

    + 1J"
(1 + 1)lti
     =  0.163 for thermal incinerator, catalytic incinerator,
            condenser, piping
                 :          '     I '  , ;         ,     ',  ' '' ,1   i
Taxes, insurance, and administration: 0.04 x Total  installed capital  cdstk
Maintenance cost:  0.05  x Total installed capital
Operating hours:  8000 hours/yr

                               5-6
-------
                           FOOTNOTES FOR Table 5-3

a
 Includes wages plus 40 percent for labor-related  administrative  and
 overhead costs.
b
 Blackburn, J.W. Control  Device Evaluation:  Thermal  Oxidation,  Report
 No. 1 in Organic Chemical  Manufacturing, Volume 4.   U.S.  Environmental
 Protection Agency.  Research Triangle Park, N.C.  Publication No.
 EPA-450/3-80-026.  December 1980.
c
 0.5 man-hours/shift x 8600 hrs/yr  -f 8 hrs/shift +  15 percent of  the
 operating labor for supervisory costs.
d
 1 man-hour/week x 8600 hrs/yr T 8  hrs/shift f 21  shifts/week + 15  percent
 of operating labor for supervisory costs.
e
 Memo from Chasko and Porter, EPA,  September 17, 1980.  Guidance  for
 developing CT6D Cost Chapters.

 Memo from Al Wehe, to Information  Analysis  Working  Group  for the
 Industrial Boiler Working Group.  April  23, 1981.   IFCAM  Modification:

     Projected 1985 price in 1978 dollars is $4.91  + $.60  delivery  charge
     per MMBtu.

     Projected 1990 price in 1978 dollars is $5.55  + $0.61 delivery charge
     per MMBtu.

     By linear interpolation between $4.91 and $5.55/MMBtu;  1988  price
     in 1978 dollars = $5.29/MMBtu.

     Using GNP implicit price deflator index:  4th  quarter 1978 of
     154.99 and 2nd quarter 1980 of 175.28;  1988 price in  1980  dollars =
     175.28/154.99 x 5.29 = $5.98/MMBtu.

     Assumed higher heating value of 1040 Btu/scf  at 16°C(60°F).
g
 Neve rill , R.B. Capital  and Operating Costs  of Selected Air Pollution
 Control Systems." U.S. Environmental  Protection Agency.   Research
 Triangle Park, N.C. Publication No. EPA-450/5-80-002.  December  1978.
 p. 3-12:

     $5.04/1000 Ib steam, 4th quartar 1977.

     Using GNP implicit price deflator index: 4th  quarter  1977  of 142.91
     and 2nd quarter 1980 of 175.28; updated steam  price = 175.28/142.91
     x $5.04 = $6.18/1000 Ib steam.
h
 Peters, M.S. and K.D. Timmerhaus.   Plant Design and Economics  for
 Chemical Engineers.  McGraw-Hill  Book Co.New York, N.Y.Third Edition.
 1980.  p. 881.

 90 percent of styrene price given  in Chemical  Marketing Reporter.

                               5-7
-------
                    FOOTNOTES FOR Table 5-3 (Concluded)

Average equipment lives given by Neverill  in reference cited in g.,
p« 3-16.
                           .    .                    ,  „             .   ,
Fugitive Emission Sources of Organic Compounds ~ Additional  Information
on Emissions, Emission Reductions, and Costs.  U.S. Environmental
Protection Agency.  Research Triangle Park, N.C. Publication No. EPA-
450/3-82-010.  April 1982. p. 5-16.

Per reference cited in footnote k:

    9 percent of total  installed capital  costs for maintenance
    and miscellaneous charges - 4 percent of total  installed capital
    costs for taxes, insurance and administration (equivalent to
    miscellaneous).
                               5-8
-------
     In order to prevent an explosion hazard and satisfy insurance
requirements, dilution air was added to any individual  or combined waste
stream with both a lower heating value between 13 and 50 Btu/scf at 0°C
(32°F) (about 25 and 100 percent of the lower explosive limit) and an
oxygen concentration of 12 percent or greater by volume.  Dilution air
was added to reduce the lower heating value of the stream to below
13 Btu/scf.  (Adding dilution air is a more conservative assumption than
the alternative of adding natural gas and is probably more realistic as
other streams often have enough heat content to sustain the combustion
of the combined stream.)
     The combustion products were then calculated assuming 18 percent
excess air for required combustion air, but 0 percent excess air for
oxygen in the waste gas, i.e., oxygen thoroughly mixed with VOC in waste
gas.  The procedure includes a calculation of auxiliary fuel requirements
for streams  (usually with heating values less than 60 Btu/scf) unable to
achieve stable combustion at 870°C (1600°F) or greater.  Natural gas was
assumed as the auxiliary fuel as it was noted by vendors as the primary
fuel now being used by industry.  Natural gas requirements were calculated
using a heat and mass balance assuming a 10 percent heat loss in the
incinerator.  Minimum auxiliary fuel requirements for low heating value
streams were set at 5 Btu/scf to ensure flame stability.4
     For streams able to maintain combustion at 870°C (1600°F), fuel was
added for  flame stability in amounts that provided as much as 13 percent
of the lower heating value of the waste gas for streams with heating
values of  650 Btu/scf or less.  For streams containing more than 650 Btu/scf,
flame stability fuel requirements were assumed to be zero since coke
oven  gas,  which sustains a stable flame, contains only about 590 Btu/scf.
In order to  prevent damage to incinerator construction materials, quench
air was added to reduce the combustion temperature to below the incinerator
design temperature of 980°C  (1800°F) for the cost curve given by IT
Enviroscience.5
      The total  flue gas was then calculated by summing the products of
combustion of the waste gas and natural gas along with the dilution air.
The required combustion chamber  volume was then calculated for  a residence
time  of 0.75 sec, conservatively oversizing by 5 percent according to
standard industry practice.^  The design procedure assumed a minimum
                                   5-9
-------
commercially available size of"1.01 m3 (35.7 ft3) based on vendor information7
and a maximum shop-assembled unit size of 205 m3 (7,238 ft3).8
     The design procedure would allow for pretreating of combustion air,
natural gas, and when permitted by insurance guidelines, waste gas using
a recuperative heat exchanger in order to reduce the natural  gas required
to maintain a 870°C (1600°F) combustion temperature.  However, all
streams to thermal incinerators costed for these polymers and resins
had sufficient waste gas heating values to combust at 870°C (1600°F)
without preheating the input streams.  If a plant had a use for it,
heat could be recovered.  (In fact, a waste heat boiler can be used to
generate steam, generally with a net cost savings.)
     5.1.1.2  Thermal Incinerator Costing.  Thermal incinerator purchase
costs were taken directly from the IT Enviroscience graph for the calculated
combustion chamber volume.5   (Essentially equivalent purchase costs
would be obtained by using data from the GARD manual.2)  A retrofit
installation cost factor of 5.29 was used based on the Enviroscience
document (see Table 5-2).9
     The installed cost of one 150-ft. duct to the  incinerator and its
associated fan and stack were also taken directly from the IT Enviroscience
study.10  A minimum cost of $70,000  (December 1979  dollars) was assumed
for waste gas  streams with flows below 500  scfm.  The costs of piping or
ducting from the  process sources to the  150-ft.  duct costed above were
estimated for  70  feet long  "source legs."11   For flows  less than  700 scfm,
an economic pipe  diameter was calculated based on an equation in the
Chemical Engineer's Handbook12 and simplified as suggested by
Chontos.13*14*15   The next  larger  size  (inner diameter)  of schedule
40 pipe was  selected  unless the  calculated  size  was within 10 percent of
the  size interval  between the next smaller  and next larger standard
 sizes.   For flows of  700  scfm and  greater,  duct  sizes  were calculated
 assuming a  velocity  of  2,000  fpm for flows  of 60,000 acfm or  less and
 5,000 fpm  for flows  greater than 60,000  acfm.   Duct sizes that  were
multiples  of 3-inches were  used.  (See  Section  E.6  for  detailed  design
 and  cost procedures  for piping  and ducting.)
      Piping costs were  based  on  those given in the Richardson Engineering
 Services Rapid Construction Estimating  Cost System16 as combined  for
 70 ft. source legs and  500  ft.  and 2,000 ft.  pipelines  for the  cost
                                   5-10
-------
analysis of the Distillation NSPS.l?  Ducting costs were calculated
based on the installed cost equations given in the GARD Manual.18
     Installed costs were put on a June 1980 basis using the following
Chemical Engineering Plant Cost Indices:  the overall  index for  thermal
incinerators; the pipes, valves, and fittings index for piping;  and the
fabricated equipment index for ducts, fans, and stacks.  Annualized
costs were calculated using the factors in Table 5-3.   The electricity
required was calculated assuming a 6-inch 1^0 pressure drop across the
system and a blower efficiency of 60 percent.
5.1.2  Flare Design and Cost Basis
     Elevated steam-assisted flares were costed based  upon 60 fps and
300 Btu/scf and standard design techniques.  Associated piping and
ducting from the process sources to a header and from  a header to the
flare were conservatively designed for costing purposes.  Operating
costs for utilities were based on industry practice.
     5.1.2.1  Flare Design.  Design of flare systems for the combinations
of waste streams was based on standard flare design equations for diameter
and height presented by IT Enviroscience.19  These equations were simplified
to functions of the following waste gas characteristics:  volumetric
flow rate, lower heating value, temperature, and molecular weight.  A
minimum commercially available diameter of 2 inches was assumed.  The
height correlation premise is design of a flare that will not generate a
lethal radiative heat level (1500 Btu/ft2 hr, including solar radiation20)
at the base of the flare (considering the effect of wind).  Heights in
5-foot multiples with a minimum of 30 ft. were used.21
     Supplemental fuel, natural gas, is added to increase the heating
value to 300 Btu/scf to help ensure 98 percent VOC destruction.   For
flares with diameters of 24-inches or less, this natural gas was assumed
to be premixed with the waste gas and to exit out the  stack.  For larger
flares, a gas ring at the flare tip was assumed because such separate
piping is more economical than increasing the flare stack size for large
diameter.
     Purge gas also may be required to prevent air intrusion and
flashback.  A purge velocity requirement of 1 fps was  assumed during
                                  5-11
-------
periods of continuous flow for standard systems without seals.22  No
purge gas is needed for either model plant under consideration.
     Natural gas consumption at a rate of 80 scfh per pilot flame to
ensure ignition and combustion was assumed.  The number of pilots was
based on diameter according to available commercial  equipment.23
     Steam was added to produce smokeless combustion through a combined
mixing and quenching effect.  A steam ring at the flare tip was used to
add steam at a rate of 0.4 Ib steam/1b of hydrocarbons (VOC plus methane
                                                     ,„• ,i,         ,     I
and ethane) in the continuous stream.24  Availability and deliverability
of this quantity of steam was assumed.
     Piping (for flows less than 700 scfm) or ducting (for flows equal
to or greater than 700 scfm) was designed from the process sources to a
header combining the streams (via "source legs") and from the header to
the base of the flare  (via "pipelines").  Since it is usual industry
practice, adequate pressure (approximately 3 to 4 psig) was assumed
available to transport all waste gas streams without use of a compressor
or fan.  The source legs were assumed to be 70 feet  in length,11 while
the length of pipelines to the flare was based on the horizontal distance
required to provide the safe radiation level for continuous working
(440 Btu/hr-ft2, including solar radiation23).  The  sizes of piping and
ducting were estimated as for thermal incinerators  (see Appendix E-6).
     5.1.2.2  Flare Costing.  Flare purchase costs; were based on costs
                        --•-•'-                       , ,           ,        ]
for diameters from 2 to 24 inches and heights  from  20 to 200 feet provided
by National AirOil Burner, Inc., (NAO) during  November 1982.23  These
costs are October 1982 prices of self-supporting flares without ladders
and platforms for heights of 40 feet  and  less  and of guyed flares with
ladders and platforms  for heights of  50 feet and greater.  Flare purchase
costs were  estimated by either choosing the value provided for  the
required height and diameter or using two  correlations developed from
the NAO data  for  purchase cost as a function of  height and diameter.
(One correlation  for heights of 40  feet and less and one for heights  of
50 feet and greater).  A  retrofit installation factor of 2.65  (see
Table 5-2)  was  used to estimate installed  flare  costs.
                                   5-12
-------
     Installed piping and ducting costs were estimated as  noted  for
thermal  incinerators (see Appendix E.6).  Installed costs  were put  on  a
June 1980 basis using the following Chemical  Engineering Plant Cost
Indices:  the overall  index for flares; the pipes,  valves, and fittings
index for piping; and the fabricated equipment index for ducting.
Annualized costs were calculated using the factors  presented in  Table  5-3.
5.1.3  Catalytic Incinerator Design and Cost Basis
     Catalytic incinerators are generally cost effective VOC control
devices for low concentration streams.  The catalyst increases the
chemical rate of oxidation allowing the reaction to proceed at a lower
energy level (temperature) and thus requiring a smaller oxidation  chamber,
less expensive materials, and much less auxiliary fuel (especially  for
low concentration streams) than required by a thermal  incinerator.   The
primary determinant of catalytic incinerator capital  cost is volumetric
flow rate.  Annual  operating costs are dependent on emission rates,
molecular weights, VOC concentration, and temperature.  Catalytic
incineration in conjunction with a recuperative heat exchanger can
reduce overall fuel requirements.
     5.1.3.1  Catalytic Incinerator Design.  The basic equipment components
of a catalytic incinerator include a blower, burner, mixing chamber,
catalyst bed, an optional heat exchanger, stack, controls, instrumentation,
and control panels.  The burner is used to preheat the gas to catalyst
temperature.  There is essentially no fume retention  requirement.   The
preheat temperature is determined by the VOC content of the combined
waste gas and combustion air, the VOC destruction efficiency, and  the
type and amount of catalyst required.  A sufficient amount of air  must
be available  in the gas or be supplied to the preheater for VOC  combustion.
(All the gas streams for which catalytic incinerator control system
costs were  developed are dilute enough  in air and therefore require no
additional  combustion air.)  The VOL components contained in the gas
streams  include ethylene, n-hexane, and other easily oxidizable  components.
These VOC components have catalytic ignition temperatures below 315°C
(600°F).  The catalyst bed outlet temperature is determined by gas VOC_
content.  Catalysts can be operated up to a temperature of 700°C (1,300°F).
However, continuous use of the catalyst at this high  temperature may
cause accelerated thermal aging due to recrystall ization.
                                  5-13
-------
     The catalyst bed size required depends upon the type of catalyst used
and the VOC destruction efficiency desired.  Heat exchanger requirements
are determined by gas inlet temperature and preheater temperature.   A
minimum practical heat exchanger efficiency is about 30 percent;  a  maximum
of 65 percent was assumed for this analysis.  Gas temperature, preheater
temperature, gas dew point temperature, and gas VOC content determine the
maximum feasible heat exchanger efficiency.  A stack is used to vent the  '
flue gas to the atmosphere.                                              '
     Fuel gas requirements were calculated based on the heat required
for a preheat temperature of 315°C (600°F), plus 10 percent for auxiliary
fuel.  The  fuel was  assumed to be natural  gas, although oil (No. 1 or 2)
can be used.  Electricity demand was based on pressure drops of 4 inches
water for systems without heat recovery and 10 incheswater for systems
with heat recovery,  a  conversion  rate  of 0.0001575  hp/in. water, 65  percent
motor efficiency, and  10  percent  additional electricity  required for
instrumentation,  controls,  and miscellaneous.  A catalyst  requirement of
2.25 ft3/l»000  scfm  was  assumed for  98 percent efficiency.25   Catalyst
replacement every three years was assumed.
     5.1.3.2  Catalytic  Incinerator  Costing.  Calculations  for capital
cost estimates  were  based on  equipment purchase  costs  obtained from
vendors  for all  basic components  and the  application  of  direct and
 indirect cost factors.25,26,27   Purchase  cost equations  were  developed
based  on vendor third quarter 1982 purchase costs  of catalyst incinerator
 systems with and without heat exchangers  for sizes from  1,000 scfm to
 50,000 scfm.  The cost data are based  on  carbon steel  material for
 incinerator systems  and stainless steel for heat exchangers.   Catalytic
 Incinerator systems  of gas volumes higher than 50,000 scfm can be
 estimated by considering two equal volume units in the system.  A  minimum
 available unit size of 500 scfm was assumed2^29;  the installed cost of
 this minimum size unit, which can be used without addition of gas  or
 air for stream flows greater than about 150 scfm29, was estimated to be
 $53,000 (June 1980).  Heat exchangers for small size systems are costly
 and may not be practical.  The direct and indirect cost component
 factors used for estimating capital costs of catalytic incinerator
 systems with no  heat  exchangers  and for heat exchangers were  presented
 in Table 5-2.   Installed costs of piping, ducts,  fans, and stacks were
                                    5-14

-------
estimated by the same procedure as for thermal  incinerators.  Installed
costs were put on a June 1980 basis using the following Chemical  Engineering
Plant Cost indicies:  the overall index for catalytic incinerators; the
pipes, valves, and fittings index for piping; and the fabricated  equipment
index for ducts, fans, and stacks.  Annualized costs were calculated
using the factors in Table 5-3.
5.1.4  Condenser Design and Cost Basis
     This section outlines the procedures used for sizing and estimating
the costs of surface condenser systems applied to the gaseous streams
from the continuous process polystyrene model plant.  Existing polystyrene
processes may emit either styrene in steam or styrene in air, depending
on the type of vacuum system used.  Styrene in steam is more readily
condensed than styrene in air and is thus less costly to control.
Design and costing were performed for both styrene-in-steam and styrene-
in-air emissions.  For styrene in steam, a condensation system was
designed that would reduce styrene emissions from 3.09 kg/1,000 kg
product and from 0.20 kg VOC/1,000 kg of product to 0.12 kg VOC/1,000 kg
of product.  Although polystyrene processes that emit styrene in  air
are expected to have emissions already around 0.12 kg/1,000 kg of product,
an analysis was performed to design a condensation system that reduced
styrene emissions in air from 3.09 kg VOC/1,000 kg product and 0.20 kg
VOC/1,000 kg of product to 0.12 kg/1,000 kg of product.  For both
design analyses, styrene emissions were assumed to be saturated in
steam (or in air) at 27°C(80°F).
     5^1.4.1  Surface Condenser Design.  The condenser system evaluated
consists of a shell and tube heat exchanger with the hot fluid in the
shell side and the cold fluid in the tube side.  The condenser system,
which condenses the vapors by isothermal condensation, is sized based on
the total heat load and the overall heat transfer coefficient which is
established from individual heat transfer coefficients of the gas stream
and the coolant.
     Total heat load was calculated using the following procedure:  the
system condensation temperature was determined from the total pressure
of the gas and vapor pressure data for styrene and steam (and styrene
in air).  As the vapor pressure data are not readily available, the
condensation temperature was estimated by trial-and-error for styrene in
                                  5-15
-------
steam and by a regression analysis of available data points3^ for
styrene in air using the Clausius Clapeyron equation which relates  the
stream pressures to the temperatures.  The total  pressure of the stream
is equal to the vapor pressures of individual  components at the condensation
                                              ' • i '••   FN ,,• 14' I  '  '".i  .i1.' ,...•,  1
temperature.  Once the condensation temperature was known, the total
heat load of the condenser was determined from the latent heat contents
of styrene and steam and, for styrene in air,  from the latent heat
content of the condensed styrene and the sensible heat changes of styrene
and air.  The coolant is selected based on the condensation temperature.
     For styrene in steam, no detailed calculations were made to determine
the individual and overall heat transfer coefficients.  Since the streams
under consideration contain low amounts of styrene, the overall heat
transfer coefficient was estimated based on published data for steam.
For styrene in air, the styrene-in-air refrigerated condenser systems
were designed according to procedures for calculating shell side3!  and
tube side32 heat transfer coefficients and according to condenser33 and
refrigerant3^'35 characteristics given primarily in the Chemical Engineers'
Handbook ana consistent with the 8-ft. long condenser with 1-inch
outside diameter tubes assumed by Enviroscience36 for cost estimation
purposes.  Then the total heat transfer area was calculated from the
known values of total heat loads and overall heat transfer coefficient
using Fourier's general equation.
     5.1.4.2  Surface Condenser Costing.  For  styrene in steam, the
heat exchanger costs for each stream were obtained from vendors.57,38,39
For styrene in air, condensation system costs  were based on  IT
Enviroscience^O as well as vendor information.  An installation factor of
1.48 (See Table 5-2) was  used to estimate  installed condenser  costs  for
condensers  of 20 ft2 or less and 2.58 for condensers  125 ft2 or greater.
No additional piping was costed  for  condensers with 20 ft2 or  less heat
transfer area because the condenser  unit is so small  (~l-2 ft.  diameter)
that it should  be  able to be  installed  adjacent to the source.
5.2  EMISSION CONTROL COSTS                                            \
     This  section  presents the  cost  estimates  of RACT emission  control
for each of the model plants.   Tables 5-4,  5-5, and 5-6  summarize the
model  plant parameters  used  in  the  cost  analysis  arid  gives the  emission
                                   5-16
-------
Table  5-4.    POLYPROPYLENE  MODEL PLANT PARAMETERS AND  EMISSION CONTROL COSTS
Parameter
Production Capacity (Gg/yr)
VOC Concentration, wt. %
Gas Flowrate, acfm
Gas Temperature, °F
Flow Time, hr/yr
Uncontrolled Emission
Factor, kg VOC/Mg Product
Total Uncontrolled VOC
Emissions, Mg/yr
Total Existing VOC
Emissions, Mg/yr«
Projected Control Device^
Total Plant3
141
47c
63.3
397
108
8,000
36. 6d
5,165
847
TI, F
PR

100
26
129
8,000
4.07
574
57
F
Process Section^
MR

100
147
133
8,000
30
4,224
422
F
PF

9.3
222
85
8,000
2.6
367
367
TI
   Assumed VOC Reduction
     Efficiency, %

   Controlled VOC
     Emissions, Mg/yr

   Installed Capital  Cost, $
   Annualized Cost, $/yr
     98
    103

635,900  (TI)
 90,600  (F)

186,700  (TI)
 65,700  (F)
                     98
    12

27,200 (F)9
21,500
25,000 (F)9
21,400 (F)h
                                     98
    85

64,200 (F)9
31,100 (F)h

53,800 (F)9
30,400
                                                       98
414,1009
377,500h

130,8009
118,50Qh
    Excludes emissions  from raw material preparation.
    D
    Parameters are provided on a process section  across lines basis and represent initial stream
    conditions from the sources.

    Assumed plant has 3 process lines, each at 47 Gg/yr production capacity.
    i
    Including stream G, dryer ventj at 0.6 kg VOC/Mg product.

    Based on assumed 90 percent control of selected streams (given in Table 4-4, footnote a).
    f
    TI = thermal  incinerator; F = flare; CI = catalytic incinerator.

    Costs for emission  control across lines (i.e., at 141 Gg capacity).
    h                                       i
    Costs for emission  control within a single line (i.e., at 47 Gg capacity).
                                            5-17
-------
     Table  5-5.   HIGH-DENSITY  POLYETHYLENE MODEL PLANT
             PARAMETERS  AND  EMISSION CONTROL COSTS
Parameter
Production Capacity (Gg/yr)
VOC Concentration, wt. S
Gas Flowate, acfra
Gas Temperature, °F
Flow Time, hr/yr
Uncontrolled Emission
Factor, kg VOC/Mg Product
Total Uncontrolled VOC
Emissions, Hg/yr
Total Existing VOC
Emissions, Mg/yr
Projected Control Device^
Total Plants
214
71. 3C
25.6
912
70
8,000
13.1
2,805*
359
TI or F
Process
MR

99
209
70
8,000
12.7
2,718
272
F
Section5
PF

0.7
703
70
8,000
0.406
87
87
CI
Assumed VOC  Reduction
  Efficiency, %

Controlled VOC
  Emissions, Hg/yr

Installed Capital Cost, $
Annual 1 zed Cost, $
     98
     56

557,400  (TI)
 54,500  (F)

166,000  (TI)
 47,400.(F)
    98
    54

57,300 (F)e
30,900 (F)?

47,960 (F)e
28,900 (F)f
     98
201,600  (CI)e
149,900  (CI)f

 71,600  (CI)e
 56,200  (CI)f
 Excludes emissions from raw material preparation.
t>
 Parameters are provided on a process section across  lines bstsis and represent
 initial  stream conditions from the sources.

 Assumed  plant has 3 process lines, each at 71.3 Gg/yr.
d
 TI » thermal Incinerator; F » flare; CI * catalytic  incinerattor.

 Costs for emission control across lines (i.e., at 214 Gg capacity).

 Costs for emission control within a single line (I.e.,  at 71.3 Gg capacity).
                                  5-18
                                                                                             '.,, itij;	
-------
Table  5-6.    POLYSTYRENE  MODEL PLANT PARAMETERS AND EMISSION  CONTROL  COSTS
Parameter
     Styrene in Steam
Total       Process Sectionb
Plants             MR
      Styrene in Air
Total      Process Section13
Planta            MR

Production Capacity (Gg/yr)
VOC Concentration, wt. %
Gas Flowrate, acfm
Gas Temperature, °F
Flow Time, hr/yr
Uncontrolled Emission
Factor, kg VOC/Mg Product
Total Uncontrolled VOC
Emissions, Mg/yr
73.5
36.75C
15. 5d
99.5
210
8,000
3.09
227

15. 5d
49.8
210
8,000
3.09
• 114
73.5
36.75C
Id
14
80.6
8,000
0.2
14.7

id
7
80.6
8,000
0.2
7.4
Total  Existing VOC
Emissions, Mg/yr
Projected Control Device
Assumed VOC Reduction
Efficiency**, %
Controlled VOC
Emissions, Mg/yr
Installed Capital Cost, $
Annuali zed Costf, $
227
Condenser
96.1
8.8
28,000
-146,700
114
Condenser
96.1
4.4
28,000
-69 ,200
14.7
Condenser
40
8.8
32,300
5,660
7.4
Condenser
40
4.4
32,300
7,735
 Excludes emissions from  raw material  storage  (stream A) and product finishing (Stream D).
a
 Parameters are provided  on a per process line basis and represent  initial stream  conditions from
 the sources.
c
 Assumed plant has 2 process lines, each at  36.75 Gg/yr.
d
 Weight % of total mass of stream.
e
 From uncontrolled emission rate to 0.12 kg  VOC/Mg product.


 Includes Styrene recovery credit.
                                             5-19
-------
                                                                            *!"!	I!"1"'!1!	IlllilSilrll,
reductions from uncontrolled to RACT levels and the installed capital
costs and annualized costs of achieving RACT for the three polymers.
The specific assumptions and breakdowns of capital  and annual costs and
recovery credits, where appropriate, are given for each model plant in
the following sections.
     Three cost analyses were performed for each polymer.  These analyses
were made in order to reflect various control costs and emission
reductions associated with applying RACT at plants that have different  ;
existing levels of control.  The actual costs and emission reductions
will depend upon the actual existing control level, stream characteristics,
                                                                        i
potential stream combination, and potential utilization of existing
control devices.
     The first  cost estimate is based on combining all continuous
streams that were judged to be reasonable to control and delivering the|
combined stream to a single control device.  Since the reduced cost of
piping will generally not offset the increased cost of multiple control
                                                                        i
equipment units, the use of a single control device for a plant is
usually the lower cost option open to a plant.
     The second cost estimate is based on  combining all continuous
streams from one type of process section in a  plant and delivering the
combined stream to a single control device.  This  analysis  reflects,
for example, a  potential situation in which a  plant may be  controlling
all other emissions except those from product  finishing and  in order  to^
control product finishing  emissions a new  control  device  is  required.
     The third  cost estimate  is, in a  sense, a subset  of  the second
                                                       i:              i
cost estimate where a  plant may be controlling,  for example, product
finishing emissions from  some  of the  lines.  This  estimate  reflects  a
"worst  case" scenario  in  which a new  control device is  required to
control the emissions  from a  single  process  section  (e.g.,  product
finishing)  in  a single  process line.
5.2.1   Polypropylene  (PP)
     The  first cost estimate  was developed for 98  percent VOC destruction
by both thermal incinerators  and flare  control of  the  combined  continuous
emission  streams from the liquid-phase polypropylene  process,  the cost
analysis  is based  on  a fluidized bed  dryer with emissions of 0.6  kg
VOC/1000  kg of product.   (Some other dryers  are potentially larger
                                   5-20
-------
emitters, while other dryers, e.g., those using recycled nitrogen,  are
extremely small emitters.)  The dryer emissions were further diluted
with air to prevent a potential explosive hazard.  The combined emission
stream from polypropylene plants is very rich in VOC so that quench
dilution air must be added in the incinerator combustion chamber to
keep the combustion chamber temperature below the limit for the con-
struction materials of 980°C (1800°F).  (Alternatively, the combined
stream is often diluted with nitrogen - about 10-30 volume percent  of
the total diluted stream - to keep the lower heating value in the
desired range of 1000-1100 Btu/scf.  This nitrogen-diluted stream has
characteristics similar to those of natural  gas and, thus, can be used
readily in boilers as a fuel supplement.) Similarly, no auxiliary fuel
is required for flaring.
     Table 5-7 summarizes the results of the cost analysis for the
polypropylene model plant.  Breakdowns of capital and operating costs
are presented for both thermal  incinerator and flare systems.  The  use
of a boiler to control VOC emissions was not costed because boiler
availability and operating practice are both site-specific; however,
costs for this control option basically consist of piping costs which
would be negligible when compared with the expected energy credit.   The
total installed capital cost of RACT is $635,800 for a thermal incinerator
system and $90,600 for a flare system.  The annualized cost is $186,700
per year for an incinerator system and $65,700 per year for a flare.
     The second cost estimate was developed based on controlling
emissions separately from each process section across process lines in
a model plant.  The third cost estimate was developed based on controlling
emissions separately from each process section within a process line.
Tables 5-8 and 5-9 summarize the results of these two additional cost
analyses.
5.2.2  High-Density Polyethylene (HDPE)
     The first cost estimate for thermal incinerators and flares achieving
98 percent destruction of VOC emissions from the high-density polyethylene
model plant was based on one stream combining the three continuous
emission streams:  ethylene recycle treaters, dryer, and continuous
mixer vents.  An air-fluidized dryer with emissions of 0.4 kg VOC/1000
kg of product was assumed.  As noted for the polypropylene model plant,
                                  5-21
-------
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other dryers may have higher or lower emissions.  The combined stream
characteristics were calculated based on the individual  stream character-
istics and compositions given in Table 2-6 and an assumed composition
of 0.7 percent isobutane in air (thus requiring no further dilution to
reduce the lower heating value to below 25 percent of the lower explosive
limit in order to prevent an explosion hazard).  Because of the substantial
VOC content of the combined waste gas stream, quench air is required to
reduce the combustion temperature of the incinerator and no auxiliary
natural gas is required for flaring.
     Table 5-10 summarizes the results of the cost analysis for the
high-density polyethylene model plant.  Breakdowns of capital  and operating
costs are presented for both thermal incinerators and flare systems.
The total installed capital cost estimated for RACT is $557,400 for a
thermal incinerator system and $54,500 for a flare system.  The annualized
RACT cost estimates are $166,000 for a thermal incinerator and $47,400
for a flare system.  As was done for the polypropylene model plant,
two additional cost analyses were undertaken.  The results of these two
analyses are summarized in Table 5-11.
5.2.3  Polystyrene (PS)
     Costs of achieving RACT for polystyrene continuous processes were
estimated based on further condensation of VOC emitted from the two
vents from the system recovering unreacted styrene monomer:  the styrene
condenser vent and the styrene recovery unit condenser vent.  The extruder
quench vent, the other stream within the scope of this CT6, contains
only a trace of styrene in steam and was not considered for control
under RACT.  The styrene emissions from the two streams were combined
and cooled to reduce gaseous emissions to 0.12 kg VOC/1000 kg of product.
Current industry control  is in a transitional period in which vacuum
pumps are replacing steam eductors to produce the required vacuum.  This
transition is taking place because of cost incentives to recover styrene
as vacuum pumps result in lower emissions of styrene to the atmosphere.
Both an "uncontrolled" emission level of 3.09 kg VOC/1000 kg of product
and an already relatively well-controlled emission level of 0.20 kg
VOC/1000 kg of product were examined in the cost basis.   The higher level
is based on one plant that is already in the process of reducing emissions to
                                  5-25
-------










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-------
below the 0.12 kg VOC/1000 kg of product level  (through  the use of
vacuum pumps) because of economic incentives.  The 0.20  kg VOC/1000 kg
of product VOC level is based on the plant that currently has the
greatest known emissions.  Both emission levels were costed for styrene-
in-steam emissions as well as for styrene-in-air emissions.
     Both of the units examined for recovery of styrene  in steam require
only a minimum commercially available size condenser unit with 20 ft2
heat transfer area.  Therefore, as shown in Table 5-12,  the only difference
in the costs of the two units is the amount of the recovery credit.
The total installed capital cost and annualized cost of both units are
$28,000 and $8,300  per year, respectively.  The net annualized costs
considering recovery credit are $-147,000 per year for a reduction from
3.09 kg/1000 kg and $4,130 per year for a reduction from 0.20 kg/1000 kg.
     For the recovery of  styrene in air, a minimum commercially available
size condenser unit with  20 ft2 heat transfer  area is required when the
uncontrolled emission  rate is 0.2 kg VOC/Mg  of product.   If  the
uncontrolled emission  rate is 3.09  kg VOC/Mg of product, then a condenser
with 185 ft2 heat transfer area is  required  to remove the  styrene  from
the  styrene-in-air  emissions.
     Only one  additional  cost  analysis  was  undertaken for  the  polystyrene
plant  because  the model  plant consists  of only two  process lines and only
one  process  section for which  RACT  is being  recommended.   The  additional
cost analysis  looks at controlling  emissions from a single process
line.   The  results  of this cost analysis  are summarized in Table 5-13.
5.3   COST EFFECTIVENESS OF RACT
      The annualized cost effectiveness  values  (net annualized  cost per
megagram of VOC  emission reduction) are given  in  Tables 5-14,  5-15,  and
 5-16 for the various control  techniques and model plant combinations.
The  estimated costs of emission reduction (using  the  lowest cost combustion
 control option of flares for PP and HOPE) are  only $13  and $17 per
 megagram of VOC reduced from uncontrolled levels, for polypropylene and
 high-density polyethylene, respectively.   Even for a reduction from the
 assumed upper level of existing control for which additional or replacement
 control might be required, the cost of the same control techniques
 would be about $88 and $160, per megagram.  For the polystyrene model
 plant, the condenser analysis results in a range in the potential  cost
                                   5-28
-------
       Table 5-12.   COST  ANALYSIS FOR POLYSTYRENE  MODEL  PLANT
Item
Installed Cost, "$
-Purchased Equipment
-Installation
Total Installed
Annual 1 zed Cost, $/yr
Direct
-Operating Labor
-Maintenance
-Natural Gas
-Electricity
-Steam
Subtotal
Indi rect
-Capital Recovery
-Tax, Insurance &
Administration
Subtotal
Styrene in Steam

11,300*
16,700a

1,080
1,400
140
5
17330"

4,550
1,120
Styrene
0.2 kg VOC
per Mg product
13,000*
19,300*
32,300''

1,080
1,620
570
17255"

5,260
1.290
6,550
in Air
3.09 kg VOC
per Mg product
41,300*
65,200*
136,000°

15,770
6,800
14,000
65d
36 ,600

22,130
5.440
27,570
  Recovery Credit
  Total  (Direct + Indirect
         Recovery Credit)
 155,0008
-146,700*
4,170
5,660
                                                                         155,000
                                                                         -90,800
 Includes only condenser and  refrigeration unit  costs.
b
 Condenser only 1-2 ft 1n diameter so no piping, etc. beyond that in  installation
 cost considered necessary.

 Includes $29,500 for piping.
i
 Cost is for make-up coolant.
a
 From uncontrolled emission rate of 3.09 kg VOC/Mg product.
f
 From uncontrolled emission rate of 0.2 kg VOC/Mg product.
                                       5-29
-------
      Table  5-13.
COST ANALYSIS FOR POLYSTYRENE
     WITHIN A PROCESS  LIME
PROCESS  SECTIONS
Item
Installed Cost, $"
-Purchased Equipment
-Installation
Total Installed
Annual! zed Cost, $/yr
Dl rect
-Operating Labor
-Maintenance
-Natural Gas
-Electricity
-Steam
Subtotal
Indirect
-Capital Recovery
-Tax, Insurance &
Administration
Subtotal
Recovery Credit
Total (Direct + Indirect -
Recovery Credit)
,
Styrene 1n Steam

ll,300a
16,700*
28,odbb

1 ,080
1,400
140
' • 5
4,550
1,120
5,670
77 ,500e
(2,090)f
-69,200e
(6,210)f
Material Recovery
Styrene
0.2 kg VOC
per Mg product
.; , '!• i 	 ' .•
13 ,,000*
19,300*
32j355b

1 .,080
1,620
isV
15d
3,270
5 ,260
1 ,290
' 6",5SO~
2,085
7,735
]
in A1r
3.09 kg VOC
per Mg product
•I,
, . " ' , i
30,900a
48,8003
94,45'OC

15,770
4,720
5,120 i
65d
25,680
15,370
3.780
19,150
77,485
-32,660
 Includes only condenser and  refrigeration unit costs.
1)
 condenser only 1-2 ft in diameter so no  piping, etc. beyond that  In installation cost
 considered necessary.
 Includes $14,740 for piping.
i                                      '                            ,          •
 Cost  Is for make-up coolants.
 From  uncontrolled emission rate of 3.09  kg VOC/Mg product.
 From  uncontrolled emission rate of 0.2 kg VOC/Mg product.
                                       5-30
-------
    Table  5-14.   COST EFFECTIVENESS OF RACT APPLIED  TO  CONTINUOUS  STREAMS
                IN  THE  POLYMERS  AND  RESINS INDUSTRY,  BY MODEL PLANT
Polymer
Polypropylene*3

High-Density
Polyethylene*3

Polystyrene


VOC Reduction
Projected From
Control Device Uncontrolled
Thermal
Incinerator
Flare
Thermal
Inci nerator
Flare
Condenser
(styrene in steam)
Condenser
(styrene in air)
5,061
5,061
2,748
2,748
218C
s.gd
218C
5.9d
, Mg/yr
From
Existing
Control3
744
744
303
303
218
5.9
218 .
5.9
Cost Effectiveness,
$/Mg VOC Reduced
Annual i zed
Cost, $/yr
186,700
65,700
166,000
47,400
-146,700
4,130
-90,800
5,660
From
From Existing
Uncontrolled Control"
37
13
60
17
-670
700
-415
960
250
88.
550
160
-670
700
-415
960
 Based on assumed 90 percent existing control  of selected  streams (given in Table 4-4, footnote a)
 for PP and HOPE and 0.20 kg VOC/Hg for PS in  order to estimate an upper end of the potential
 cost effectiveness range.
b
 Boilers can also be used to achieve 98 percent VOC reduction efficiency.

 From uncontrolled emission rate of 3.09 kg VOC/Mg product.
d
 From uncontrolled emission rate of 0.2 kg VOC/Mg product.
                                         5-31
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styrene in steam reduced from 3.09 kg/Mg) to a cost of $960/Mg (styrene
1n air reduced from 0.20 kg/Mg).
                                   5-34
-------
 5.4 REFERENCES FOR CHAPTER 5

 1.  Polymer Manufacturing Industry - Background Information for Proposed
    Standards.  U.S. Environmental Protection Agency, Research Triangle
    Park,  North Carolina.  Draft EIS.  September 1983.

 2.  Neverill, R.B. Capital and Operating Costs of Selected Air Pollution
    Control Systems.  U.S. Environmental Protection Agency.  Research
    Triangle Park, N.C.  Publication No. EPA-450/5-80-002.  December 1978.
    p.  3-7 and 3-8.

 3.  Memo  from Mascone, D.C., EPA, to Farmer, J.R., EPA.  June 11, 1980.
    Thermal incinerator  performance for NSPS.

 4.  Blackburn, J.W.  Control Device Evaluation: Thermal Oxidation.  In:
    Chemical Manufacturing Volume 4:  Combustion Control Devices.  U.S.
    Environmental  Protection Agency, Research Triangle Park, N.C.
    Publication No.  EPA-450/3-80-026, December 1980.  Fig  III-2, p. III-8.

 5.  Reference 4,  Fig. A-l, p. A-3

 6.  Air Oxidation Processes in Synthetic Organic Chemical  Manufacturing
     Industry - Background  Information for  Proposed Standards.  U.S.
    Environmental  Protection Agency, Research Triangle Park, N.C.
    Draft EIS.  August 1981.  p. 8-9.

 7.   EEA.  Distillation NSPS Thermal  Incinerator Costing Computer Program
     (DSINCIN).  May  1981.  p. 4.

 8.   Reference 4,  p.  1-2.

 9.  Reference 6,  p.  G-3  and G-4.

10.   Reference 4,  Fig. V-15, curve  3, p. V-18.

11.   Memo from Senyk, David, EEA,  to  EB/S Files.  September 17, 1981.
     Piping and  compressor  cost  and  annualized cost parameters used in
     the determination'of compliance  costs  for the EB/S  industry.

12.   Perry, R.H.  and  C.H. Chilton,  eds.  Chemical Engineers'  Handbook,
     fifth edition.  New  York,  McGraw-Hill  Book Company.   1973. p. 5-31.

13.   Chontos,  L.W. Find  Economic  Pi^ Diameter via  Improved Formula.
     Chemical  Engineering.   87(12):139-142.  June 16,  1980.

14.   Memo from Desai, Tarun,  EEA,  to EB/S Files.  March  16, 1982.
     Procedure to  estimate piping  costs.

15.   Memo from Kawecki,  Tom,  EEA,  to SOCMI  Distillation  File.   November 13,
     1981.  Distillation  pipeline  costing model documentation.

16.   Richardson  Engineering Services.  Process  Plant  Construction  Cost
     Estimating  Standards, 1980-1981.  1980.

                                  5-35
-------
17   EEA.  Distillation NSPS Pipeline Costing  Computer  Program  (DMPIPE),
     1981.                                                          I

18.  Reference 2, Section 4.2, p.  4-15 through 4-28.

19.  Kalceyic, V. Control Device Evaluation: Flares and the  Use  of
     Emissions as Fuels.  In: Organic Chemical Manufacturing Volume  4:
     Combustion Control Devices.  U.S. Environmental  Protection  Agency.
     Research Triangle Park, N.C.   Publication No.  EPA-450/3-80-026.
     December 1980.

20.  Reference 19, p. IV-4.

21.  Memo from Sarausa, A.I., Energy and Environmental  Analysis, Inc.
     (EEA), to Polymers and Resins File.  May  12,,  1982.  Flare  costing
     program (FLACOS).

22.  Telecon.  Siebert Paul, PES,  with Keller, Mike,  John Zink,  Co.
     August 13, 1982.  Clarification of comments on draft polymers and
     resins CTG document.                                          ..'•

23.  Telecon. Siebert, Paul, PES with Fowler,  Ed,,  NAO.   November 5,
     1982.  Purchase costs and operating requirements of elevated flares.

24.  Telecon. Siebert, Paul, PES with Fowler,  Ed,  NAO.   November 17,
     1982.  Purchase costs and operating requirements of elevated flares.
                                                                    i
25.  Telecon. Katari, Vishnu, Pacific Environmental Services, Inc.   with
     Tucker, Larry, Met-Pro Systems Division.   October 19, 1982.  Catalytic
     incinerator system cost estimates.

26.  Telecon.  Katari, Vishnu, Pacific Environmental  Services,  Inc., with
     Kroehling, John, DuPont, Torvex Catalytic Reactor Company.   October 19,
     1982.  Catalytic incinerator systsem cost estimates.

27.  Letter from Kroehling, John, DuPont, Torvex Catalytic Reactor
     Company, to Katari, V., PES.  October 19, 182.  Catalytic  incinerator
     system cost estimates.                                         ;

28.  Key, J.A. Control Device Evaluation: Catalytic Oxidation.   In:  Chemical
     Manufacturing Volume 4: Combustion Control Devices.  U.S.  Environmental
     Protection Agency, Research Triangle Park, N.C.  Publication No. EPA-
     450/3-80-026.  December 1980.                                  '

29.  Telecon. Siebert, Paul, Pacific Environmental Services, Inc., with
     Kenson, Robert, Met-Pro Corporation, Systems Division.  July 22, 1983.
     Miminum size  catalytic  incinerator units.                      '.

30.  Reference 12, p.  3-59.
                                                                          I:."	r!" *!S*:iS,
                         5-36
-------
31.  Reference 12, pp. 10-25 through 10-28.

32.  Reference 12, pp. 10-12 through 10-15.

33.  Reference 12, pp. 11-1 through 11-18.

34.  Reference 12, pp. 3-191, 3-212 through  3-214,  and 12-46
     through 12-48.

35.  Weast, R.C., ed. Handbook of Chemistry  and Physics,
     fifty-third edition.  Cleveland, The Chemical  Rubber Company.
     1972. p. E-26.

36.  Erikson, D.G. Control Device Evaluation:  Condensation.
     In: Organic Chemical Manufacturing Volume 5: Adsorption,
     Condensation, and Absorption Devices.   U.S. Environmental
     Protection Agency, Research Triangle Park, N.C.
     Publication No.  EPA-450/3-80-027.  December 1980. p. A-3.

37.  Telecon.  Katari, Vishnu, Pacific Environmental  Services,  Inc.,
     with Mr. Ruck, Graham Company.  September 29,  1982.   Heat  exchanger
     system cost estimates.

38.  Telecon.  Katari, Vishnu, Pacific Environmental  Services,  Inc.,
     with Glower, Dove, Adams Brothers, a representative  of Graham
     Company.  September 30, 1982.  Heat exchanger  system cost  estimates.

39.  Telecon.  Katari, Vishnu, Pacific Environmental  Services,  Inc.,
     with Mahan, Randy, Brown Fintube Company.  October 7, 1982.   Heat
     exchanger system cost estimates.

40.  Reference 36, pp. A-4 and A-5.
                                  5-37
-------
-------
    APPENDIX A



LIST OF COMMENTERS
        A-l
-------
                               APPENDIX  A
                           LIST OF  COMMENTERSe
     Comment No.          Comment Date        Commenter

                                             Monsanto Company
                                             (W.G.  Bir and C.D.  Mai loch)

                                             Texas  Chemical  Council
                                             (A.H.  Nickolaus)

                                             Chemical Manufacturers
                                               Association
                                             (Geraldine V. Cox)

                                             Gulf Oil Chemicals
                                               Company
                                             (J.R.  Strausser)

                                             Polysar Inc., Resins
                                               Division
                                             (F.J.  Mitrano)

          6                                  DuPont
                                             (G. Madden)


aOnly comments on the May 1982 draft CTG document are included.
Comment Date

June 16, 1982


June 18, 1982


June 21, 1982



June 21, 1982



July 19, 1982
                                   A-2
-------
  APPENDIX B



 COMMENTS ON



MAY 1982 DRAFT



 CTG DOCUMENT
      B-l
-------
Monsanto Company
800 N. Lindbergh Boulivard
St. Louis. Missouri 63166
Phone: (3t4) 894-1000
                                June 16, 1982
                                                      Mail Zone G3WG
 Chemicals & Petroleum Branch (MD-13)             ,  IN DUPLICATE
 Emission Standard & Engineering Division
 U.S. Environmental Protection Agency
 Research Triangle Park, North Carolina 27711

 ATTENTION:  Mr. Jack R. Farmer
,                            ,              ,                , •  :

 RE:  COMMENTS ON EPA'S DRAFT CTG ENTITLED  CONTROL  OF VOLATILE ORGANIC
      COMPOUND EMISSIONS FROM MANUFACTURE OF HIGH-DENSITY POLYETHYLENE,
      POLYPROPLENE, POLYSTYRENE RESINS, MAY 1982

 Dear Mr. Farmer:
DRAFT CTG
0.11
2.96
0.133
0.15
3.35
••.;
-------
Mr. J. R. Farmer
EPA
Page 2
June 16, 1982
         As noted, there is a 20-30 fold difference between the EPA and
         CMA numbers shown.  For the model plant in the draft CTG of
         73.5 Gg capacity, the CMA emissions would run from approximately
         9-11 Mg/Yr. compared with the 246 Mg/Yr. emissions shown in the
         draft CTG.  No references were contained in the draft CTG which
         allows backtracking to the EPA basis for the quoted emissions
         factors.  Since actual experience indicates that the CMA emission
         factors are appropriate, the emissions expected from existing
         continuous polystyrene facilities are insignificant and as such, there
         is no significant nonoccupational exposure.  Monsanto contends the
         CTG is not necessary for this industry segment.

         Thermal Incineration is not an Appropriate Control..Device

         If EPA persists,  and issues a CTG for VOC emissions from polystyrene.
         units, then Monsanto disagrees with the selection of thermal incinera-
         tion as the emission control device to use (see draft CTG on page 5-1,
         where EPA states  "Thermal incinerators are the only control device
         evaluated.")

         Incinerators are  not cost-effective control devices for control of
         the insignificant VOC levels which emit from existing continuous
         polystyrene units.  As EPA stated on page 5-14 of the draft CTG
         "For each model plant, the resulting combined stream was smaller than
         the capacity of the smallest off-the-shelf incinerator available."
         Using EPA's cost  numbers of approximately $70,000 for direct and
         indirect costs(see the draft CTG page 5-15, Table 5-7), and applying
         CMA emission levels, the cost-effectiveness would run .from about
         $6400 to $7800/Mg. VOC removed (as compared to EPA's number of
         $320/Mg. VOC).

         The cost-effectiveness levels would be even higher than this if a detailed
         cost estimate were done taking into consideration factors such as:

         1.   Due to the  insignificance of the VOC stream size,  and the size of
             the incinerator, auxiliary fuel would be needed to sustain burning,
             hence, added  operating cost.

         2.   The technology of compressing the styrene monomer vapors and trans-
             porting them  for up to 1,000 ft., would promote polymerization in
             the pipe and  hence buildup which would need to be removed periodically,
             This would  also add to the operating cost.

         As  such, incineration is not an appropriate control device to use
         on  the insignificant VOC emissions which emit from existing continuous
         polystyrene units.  In addition, Monsanto strongly objects to the use
         of  its acrylonitrile incinerator data contained in Monsanto's submission
         to  EPA on November 8,  1979 (see reference 5 on draft CTG page A-27)
         as  being an equivalent technology base for styrene.   The AN data was
                                    B-3
-------
Mr. J. R. Farmer
EPA
Page 3
June 16, 1982
         submitted for another purpose and not to be extrapolated for
         polystyrene use due to the totally different chemical and
         physical properties of the two substances.

     •   Miscellaneous Items

         1.  On page 5-12, first sentence in Section 5.3.4, EPA makes the
             statement "The continous process emits unreacted styrene
             mbnomer (VOC) because the polymerization process approaches
             equilibrium before reaction completion."  This is an incorrect
             statement since the process time for a reactor system to
             reach equilibrium from the initial startup operation has no
             Impact on this CT6.  The continuous process emits unreacted
             styrene monomer (VOC) because the fresh and recycled styrene
             feed streams are partially converted to polymers in the reactors
             (i.e.-greater than 60% styrene conversion).  The amount of
             styrene converted is set by the polymer molecular weight, the
             reactor space-time-yield, and by conventional process design
             variables.  EPA should correct the first sentence in the
             referenced section accordingly.

         2.  Table 2-7 on the draft CT6 page 2-20 shows Monsanto's Long
             Beach, California plant on the list of polystyrene producers
             in nonattainment areas.  Monsanto has shut this operation down
             and hence requests that it be removed from the list of polystyrene
             producers.
                  1         '        [                   ', , • ,  i,  !   :.<,.•  ,, f    'I •	: >'-:::! '-, i;-

As documented above, Monsanto strongly encourages EPA to cease work on developing
a. CTG for existing continuous polystyrene unit, since the resulting VOC emissions
from these units are insignificant in quantity.  Monsanto would welcome the
opportunity to discuss its point further with EPA if it is necessary for further
clarification of the points above.
                                      Sincerely,
                                      W. G. Bir
                                      Engineering Group Consultant
                                      Corporate Engineering Department
                                        L
CDMrvre
                                      C. D. Malloch
                                      Regulatory Management Director, Air
                                      Environmental Policy Staff
                                   B-4
                                                            ''I'.;	:iu ,!.„ liiilli \A\ .liil'l' '< I
                                                                         ,,,'li , I, • n'hl; |i I IP II
-------
            CHEMICAL MANUFACTURERS
                                             Attachment to Oune 16, 1982 Tetter
                                             from C.D. MaTloch, Monsanto to
                                                       ~PA.
                                 October 19, 1981
Mr.  Edwin J.  Vincent
Lead Engineer
Chemical Applications Section
Chemicals and Petroleum Branch                        .  •
Office  of Air Quality Planning and Standards
United  States Environmental Protection Agency
Research Triangle Park, North Carolina  27711

     REs  CMA Comments on SPA Model Plant Emissions  Factors for
          Polymers/Resins Manufacture

Dear Mr. Vincent:

     Our Polymers/Resins Work Group has reviewed the draft document
prepared as part of the Agency's effort to  develop  an NSPS for Poly-
mers/Resins Manufacture.  The Group has focused its attention pri-
marily  on the accuracy of the emissions factors and their location
within  individual processes as based upon its  collective experience.

     As  you may know, CMA is a nonprofit trade  association made up
of approximately 184 member companies in the United States repre-
senting more than 90 percent of the domestic production capacity
for basic industrial chemicals.  CMA member companies have a direct
and critical interest in ensuring that EPA  develops emission stand-
ards when a demonstrated need is presented, that are scientifically
and technically sound, reasonable, procedurally workable, cost effec-
tive, and clearly authorized by the Clean Air  Act.  . Many of our mem-
ber companies produce Polymer and Resin products and may be impacted
by any regulations which may be based on the subject document.

     While our comments provide what we believe are  improvements to
the model plant emissions factors, we do believe that the high de-
gree of variability between individual processes used to manufacture
. the same product demonstrates the limited usefulness of the model
plant concept.  This point will be e-xcr^plified in our discussion of
 the individual emissions factors  for the specific products/processes.

 I.  Polypropylene - continuous slurry,  liquid  phase process

     A.   There are two liquid phase processes  now in commercial
         use? the large particle slurry  process and the solution
         process.  Most new  liquid phase plants employ or will
         employ the solution process.  A sizeable number but not
         all of these new plants are using or will use the high
         yield catalyst technology.  As  a result, there are some
       • solution processes  that still require  catalyst de-ashing

                              B-5            '
          Formally Manufacturing Chemists Association— Serving the Chemical industry Sine* 1872.
        25O1 M Strait. WW . Washington, DC 20037 » 7*teonen« 202/867-1100  * Telex S9617 (CMA VYSH)
-------
                                  -2-
          and removal.  Most  slurry  process  units  and low yield
          catalyst-plants employ  jacketed, continuous stirred-tank
          reactors rather than  loop  reactors.   The loop reactor  is
          more prevalent in the high yield catalyst plants.

      B.  Some slurry processes operate  at pressures as high as
          300 psig which is much1  higher  than the value cited in
          the draft.

      C.  The atactic generation  rate  of 30  percent of capacity
          cited in the draft  would be  uneconomic,   Most liquid
          phase processes have  average atactic  generation rates
          in the range of 2 percent  to 4 percent of nameplate
          with 5 percent an upper bound.

      D.  The model plant did not provide for a VOC emissions vent
          from the extrusion/pelletizing section.   Significant
          quantities of hydrocarbon  still remain in the polypro-
          pylene powder as it exits  the  dryer and  enters the"ex-
          truder feed chute.  At  this  point, the powder is in
          equilibrium  with a vapor  that can contain up to 25 per-
          cent hydrocarbon  (wt./wt.).  As a  result, there is some
          hydrocarbon loss through the extrucer/pelletizer section
          and the powder/pellet transfer system downstream of the
          product dryer.  One manufacturer has  provided an estimate
          of 2 kg VOC/1000 kg  product for this section of the pro-
          cess.          '   i      '        '    ,;;./ ;   _  ,  '  i,;,'  ';"",. .:.,„_,,
                               i           . •-, ' '"is '!,  : ;•'••"• >' ?'>;. I. rA^".'"; Vi:*^
      E.  Polypropylene units are subject to plugging, especially
          in the polymer handling sections cf the'process.  For  this
          reasqh, most units  are  provided with  emergency relief
          valves, where applicable,  throughout  the process and not
          just in the polymerization and atactic recovery sections.
          In the vast majority  of cases,  these  relief valves' are"'
          tied to "the flare header.

      Polyethylene - all products/processes
      A.
There are a considerable number  of LDPE  and  HpPE .plants
located near or integrated with  olefins  manufacturing
                                                              ng
          operations.  Some of  these  units  do  not  have  recycle
          treaters since monomer  recovery  and  purification* is accom-
          plished by recycle through  the" olefir.s manufacturing unit.
          In these cases, overall process  VOC  emissions from these
          units can be expected to be 3-10  percent lower for the LZPE
          processes and by up to  98 percent lower  for the HDPE ore-
          cesses.

III.  Low Density Polyethylene  -  all  process

      A.
                               B-6
-------
B
C.
         The model plant does not include an emissions factor  for
         the wax blovdown system.  This section of the process can
         be a source of significant ethylene losses.  The emissions
         factor is highly dependent upon the desion of the wax blow-
         down and discharge system.

         For the  liquid phase 'process, the estimated frequency for
         emergency reactor conditions  is too high.  The  assumption
         that two out of four reactors would simultaneously exoeri-
        . @ne@ an upset is extremely unlikely.  Furthermore the"esti-
         mated 130 reactor upsets per  year is atypical of industrial
         experience.

IV.  High Density Polyethylene - all process

     A.  Many plants use air-fluidi2sd rather than inert gas dryers.
         In these plants there are some VOC emissions, in'the  range
         of 0.2-0.4 kg VOC/1000 kg product.  In plants without re-
         cycle treaters, these are the major vents.  Most of these
         emissions consist of process  diluent.

     3.  Most plants use separate recycle treaters for each individual
         hydrocarbon component since they are usually recovered by
         fractional_distillation.  Therefore/ an HDPS plant recycling
         ethylene, isobutane and butane would under most circumstances
         have three treaters and the vent composition of each  treater
         would contain 100 percent of  hydrocarbon treated.

 V.  Polystyrene - bateh process

     A.  The emissions factors cited for the model plant may approxi-
         mate the average for the industry but may not adequately
         describe the emissions for the purposes cf regulation.  Batch
         plants are well suited for use in the manufacture of  a wide
         variety of'products.  Emissions factors fcr the vents for  the
         process also vary widely with the higher emissions factors
\        more likely during the manufacture of lower molecular weicht
         products.  Typical emissions  factors span the following
         ranges:

              Styrene Condenser Vent   0.25 - 0.75 kg VOC/1000  ka resir.
              Extruder Quench Ven-L     J.1S - 0.30 kg VOC/1000  kg resir.

              Reactor Heading Vent     0.15 - 1.35 kg VOC/1000  kg resin

         It is important to note that  the emissions factor for anv
         given process train will chance with product grade.

VI.  Polystyrene - continuous process

     A.  Like the batch polystyrene plants, industry's experience with
         continuous po'lystyrene plants indicates a"wide  ranee  cf emis-
         sions factors but for different reasons, 'individual  continu-
         ous process trains tend to run blocked-ir. en one' ooivmer
         grade or family of grade with relatively small  variations  i-
                             B-7
-------
        emissions factors.   However, different process trains
        dedicated to widely differing polymers may have signifi-
       • cant differences in their emissions factors.

    B.   Recent process improvements have included a shift to the
        use of vacuum pumps to generate process vacuums.  These
        pumps consume less  energy than steam eductors and also have
        lower emissions factors-.  Industry's experience with the
        use of vacuum pumps are substantially different than one
        would expect when reviewing the model plant emissions fac-
        tors.  Our experience indicates that the foliowing emissions
        factors are more typical of newer continuous polystyrene
        process technology:

             Tankage

             Styrene Condenser Venr

             Styrene Recovery Unit
               Condenser Vent
             Extruder Quench Vent    0.009-0.01 kg VOC/1000 kg resin
                                             "    ,„!  :       . '• •  j' ' "
    We trust that these comments will be considered as you revise the
draft model plant emissions factors.  We thank you  for giving us the
opportunity to review this material and are willing to meet with you
to discuss our concerns in greater  detail.  Rich Symuieski/  the Work
Group leader will contact you for follow-up in this regard.

                                Sincerely,
                                     0.01 - 0.02 kg VOC/1000 k£ resin

                                     0.03 - 0.06 kg VOC/1000 kg resin

                                     0.05 - 0.06 kg VOC/1000 kg resin
                                Janet S. Matey
                                Manager
                                Air Programs
JSM/sl
                                                                  in"  4 'it.*''.'
                              B-8
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               TEXAS CHEMIttU. 420lifli0§L
1000 BRAZOS, SUITE 200, AUSTIN, TEXAS 78701 -2476, (512) 477-4465
                                           June 18, 1982
   Mr.  Jack R.  Farmer, Chief (2)
   Chemicals and Petroleum Branch
   Emission Standards and Engineering Division (MD-13)
   Office of Air Quality Planning and Standards
   U.  S.  Environmental Protection Agency
   Research Triangle Park, North. Carolina  27711
   RE:
TCC Review Comments on the Draft CTG for Control of VOC
front Polymer/Resin. Manufacture 47FR19580 (May 6,1982)
   Dear Mr.  Farmers    .      .                          .

             The Texas. Chemical Council (TCC)  submits the attached
   comments  on the subject draft Control Technique Guideline for the
   C*n^°i ^f v?latile organic compound emissions from the manufacture
   of  high-density polyethylener polypropylene,  and polystyrene.

             Should the agency have any questions or wish to discuss
                   5* r*^5 by Our comraents, you may contact me at
                  Ext. 1277, or write our Austin office
                                      Very truly yours,
                                      A.  H. Nickolaus
                                      Chairman,  CTG Subcommittee
                                      Texas Chemical Council
  cc:  TCC Air Policy Committee
       J. B. Cox  - Exxon
       J. S. Matey - CMA
       P. J- Sienknecht - Dow
       TCC Files
  AHN/cgh
                                B-9
-------
              COMMENTS BY THE TEXAS  CHEMICAL  COUNCIL
     ON THE MAY, 1982 DRAFT CONTROL  TECHNIQUE GUIDELINE  (CTG)
     FOR CONTROL OF VOLATILE ORGANIC COMPOUND (VOC) EMISSIONS
        FROM THE MANUFACTURE OF HIGH-DENSITY  POLYETHYLENE,
     	POLYPROPYLENE, AND POLYSTYRENE
fl* ^m-,~ ?       ? Chemical Council  (TCC)  is  an  association of
85 chemical companies having more  than  67,000  employees in Texas
f?d ffPf Denting approximately 90% of the chemical industry in
the State.  Many of the polyethylene  and polypropylene plants
covered by the CTG are in Texas  and,  thus,  the proposed guide-
lines are of concern to us.

     A.   Review of Previous Comments

          This draft CTG is very similar to the April, 1981 version  •
          reviewed at the June 2-3, 1981 National Air Pollutant
          Control_Technique Advisory  Committee meeting and we are
          disappointed that it does not more fully reflect the
          TCC comments, submitted to the EPA than  (Ref. 1)..  Most
          of our previous comments-are still pertinent and are
          summarized below as they relate to -the present CTG.

          1-   The CTG Does Not Fulfill It's Stated Purpose

               The TCC continues to believe that the omission of
            •   absorption.,, and other pollutant, recovery techniques,
               and the definition,of Reasonably Available Control
             ...Technology (RACE)  exclusively in terms of thermal
               incineration is not very useful in helping the states
               proceed with their own assessment of RACT - the  •
               guideline's stated purpose.The reasoning used in
               Section 3-1 to dismiss these other technologies
             ,  presents no data and is largely specious.

          2*   RACT Should Allow Several Technologies

               In our: May 29,  1981 comments we set forth what we
               thought were excellent reasons why RACT should allow
               several, abatement technologies..  We still think they
            .  are  valid and that a- 98% reduction requirement is
               unduly stringent for RACT when compared to New Source
               Performance Standard (NSPS)  requirements and the
               levei of regulation on mobile and other sources.
               Restating the RACT recommendation to a 98%  reduction
               in. Section 4.1 of the present document from thermal
               incineration (under conditions to give a 98% reduction)
               in the April,  1981 version does not really address
               our  concern.
                              BrlO
                               \
                                             .uliili!:;	1','fJlI ,!
                                                                 ijilli"!,!"-!	iV: iS'Slli	II!
-------
                     ..- 2  -
4.
                              A 98% Abatement Standard
                        However from subsequent consul-
     be based on demonstrated levels in

SACT Should^lnclude the Use of ?!»,.».


    2s
    •    Were the data measured?
                   B-n
-------
                  - 3 -
  If not,  we think they should not be included in the
  summary.

  This CTG (p.  5-17,18)  lists four flare studies;
  Du Pont, National Aijr oil Burner, Union Carbide,
  and Seigel.  Reference 4 also lists four flare
  studies; Du Pont, Union Carbide, Seigel,  and Zink/
  Battelle.  Why were the Zink/Battelle studies which
  showed high destruction efficiencies and which were
  run for  the EPA left, out of this study?  And why was
  the National Air Oil Burner report left out of
  Reference 4?
                                 ,                  |

  The CTG's discussion of flares states that  "The uncer-
  tainity  associated with flare combustion contrasts
  starkly  with our knowledge of incinerators  and boilers.
  Evidence- to show the thoroughness of combustion effi-.  •
  ciency in these devices is ponderous."  One can be
  ignorant of anything if they refuse to study it and the
  data presented in, Appendix A is hardly ponderous.   Six
  plant scale test results are presented and  these show
  destruction efficiencies ranging from-70.3  to 99.9%.
  If the EPA applied the same critical criteria to incin-
  erators  as they do to flares they would, have to con-
  clude incinerator efficiencies ar« significantly less
  than 93%.  Based on. information from in-house experts
  wer think,all combus-tion devices will give high destrue-
  efficiencies if the pollutant does? not by-pass and
  actually experiences the flame.  Thus even  new incin-
  erators  can give poor results such as the Petro-tex
  data, if  by-passing, occurs.                   .
 The TCCT does not want to belabor this  issue but we
. -think it -is time the EPA judged data for  flares by
 the- same criteria as they use  for boilers and  incin-
 erators.. '  We realize flare efficiency  is  not readily
 measured but the test of a control technique should
 be its cost effectiveness and  efficiency,, not  its
 ease  of enforcement.
 Finally,, as the;
 by the Chemical
 underway now at.
 this; study will,
 included in any;
 efficiency.
EPA is probably aware, flare tests
Manufacturers Association should be
the John Zink plant.  Results from
be available shortly and should be
final appraisal of flare destruction
                    B-12
-------
                          - 4 -
          Emission Reductions in the April , 1981 CTG Were
          Overstated                  *
          d af t
                                               since projections

                                        * removed f rom tnis
B.   Processes Covered by the CTg
     EPA for1thS?/h^tHCie f?™ <•}»«*> Process is used by the
     5 JSJSS »* ^gn-densxty polyethylene model plant.  It

     iith S? SK ad«£° • rePresent aH other liquid-phaSe processes
     with high-efficiency catalyst that do not require cat alvst
     removal.-  But Dow, Du Pont,  and perhaps o32£ use sSlu?ion

                  Processes that are completely different ?rom  "

     co   a^0638' the ^ssion/are different? and Sey
     come out at different places .  We presume these are not
     covered by this CTG since the EPA has male nl^atualfln

                    C0ntro1 tehcnology or abatement costs for



                  Secti011 «-l be revised to make clear that the

                  ?"**"*0* apply t0 ^H-density pSyeSylene
                 , low-pressure,  slurry, liquid-phLrprocesses?


     RftCT RecoiamendatioB and Costs


                 aXjre^av Discussed the TCC believes the 98%
               recoannendation is too stringent and that a
                                                ,             .


                already have some control facilities.
                c      is not ^fleeted in the cost analysis-
                5. .  Xt should be  since the guidelinrapplies
    to .exis.tin, processes..   Some data for high-deSJiS poly
    • ' ' Plant** •' "  ,  , '


    EPS Model Plant  ,

      — Ofccoatrolled;
      "      Reduction
.Emission Factor*
   CKg/lOOQ Kg)
     121..56
       0.25
% Reduction
From Model
  Plant
           E-Z-
                                7.5
                          90
                          40
                        8-13
-------
                      - 5 -
    Plant**

     H-3
     H-4  •
     H-5
     H-6
     H-7
     H-8
Emission Factor*
  (Kg/1000 Kg)	

      2.4
      100
        1
       10
       21
       31
% Reduction
Prom Model
  Plant
"'"	:  	{'	
     81
   (696)
     92
     20
    (67)
   (147)
                               „ •• ' , 	ii, , ; • ,;,... I,	••,!',' '.| , , "  ,:11"1!'!
 * Includes fugitive and miscellaneous emissions.
   Plants include solution processes but rhis was a blind
   inquiry so specific process or producer was not connected
   to a given emission factor.

 Consider the problem of plant H-l, 3, and 5.  To reduce
 their emissions to 0.25 Kg/1000 Kg they roust now install
 an incinerator.  For plant H-5, at model plant rates, this
 will amount to a reduction of 160 Mg/yr. at a cost of
 $121,000 per year (Table 5-4) or $756/Mg, a figure over
 an order- of magnitude higher than  the $46/Mg shown in
 the CIS cost analysis .(Table 5-6).
BPA*g Attempted Justification for Requiring Incineration
or  Streams Now Being Flared    :  '	

•In.  Section. S.4.2 the EPA attempt to justify requiring
incineration, of streams now being flared-  Their whole
case rests on the assumption of a low efficiency (90%)'
for -flares and a "state-of-the-art" efficiency (98%)
for- incinerators.  As discussed earlier (A 3 and 4),
the TCC doubts that there is any significant difference
between the destructipn efficiency of flares and incin-
erators.   Without definitive data to Quantify a difference
between the two the EPA's proposal to require the replace-
ment, of existing flares with incinerators is unconscionable.
 '              •         .                             i
                                          i	,
                          The Texas... Chemical .Council
                          1000 Brazos .[Suite 200
                          Austin,  Texas  78701
                          June 18,. 1982
                    B-14.
-------
                       REFERENCES
1.
2.
Texas Chemical Council to Don R. Goodwin, EPA,  "Comments
  by the Texas Chemical Council on the Draft Control
  Technique Guideline ...,» May 29, 1981.    v-oncroi



TCC to Jack R. Farmer, EPA, "Proposed NSPS for Air-
  Oxidation Processes, Incinerator Efficiencv""; —
  January 4, 1982.                          * '
3*
               ?mission Sources of Organic Compounds -

             ^?f?f^tion On Missions, Emission Reduction
        ,. April 1982, EPA-450/3-82-010, Page 4-68
                        B-15
-------
               CHEMICAL MANUFACTURERS ASSOCIATION
G==*A--::r.= v, cox, P.I.D,
V,ce Prts'rsn:
                                        June 21, 1982
 Xr.  Jack ?..  Farmer, Chief
 Chemicals and Petroleum Branch
 Emissions Standards and Engineering Division
 Office of Air Quality Planning and Standards
 0.  S. Environmental Protection Agency                                    ;
 Research Triangle--Park, North Carolina  27711

     HE:  CMA Review Comments on the Draft  Guideline  for the Control of
          Volatile Organic Emissions from Manufacture of High Density
          Polyethylene, Polypropylene  and Polystyrene.' Resins.	

 Dear Mr. Farmer:                                                         ;

     Th-  ChezPcal Manufacturers"Association's CCMA)  Polymers and Resins Work   '
 G-OUD has" reviewed the Draft CTG for the Control of Volatile Organic Compound
 Emissions from the Manufacture of High Density Polyethylene, Polypropylene and
 Polystyrene Resins.   Our Work Group has been tracking this  effort since the
 d-velocment of Preliminary Draft documentation by the Agency and Provided
 comments on the CTG at the June 2, 1981, NAPCTAC meeting.   In general, EPA has
 not responded to our earlier comments in preparing  the  current Draft CTG.

      As you may know, CMA is a nonprofit trade association  whose company;mimbers
  represent more than 90% of the productive capacity  of  basic industrial chemicals
  within this country.

      Our comments are  focused on  the  major issues  that  must be "solved to provide
  an elective,  flexible  CTG  of  value  to  the  states  in achieving attainment with
  Se ozone  NAAQS.  Since .the Agency has  not addressed our earlier comments, they
  are still  relevant with respect  to  the  Draft CTG.

      The issues addressed in today's  comments include:      •

      o   T'-e definition of RACT is more, typical of LAER  and  is inconsistent
          with RACT  levels defined by other CTG's for VOC emissions  reduction.
          tion Units where flares are  allowed.

          The model olant description  and emissions.factors for polystyrene
          manufacture are not  representative of -current industry Practice and
          overstate both the emissions and the cost effectiveness of RACT con-
          trol for these .sources.         B-16
              Formerly Manufacturing Chemists Associiiion-Serving the Chemical Industry Since 1872,
             253-, M Street. NW • Washington. DC 20037 • Telephone 202/887-1260 • Tele> S9617 (CMA WSH)
                                .                      	   •          '|
-------
• Mr. Jack R. Farmer
 June 21, 1982
 Page Two
      o  The Agency has relied on obsolete information in develooina incinerator
         costs which overstates the cost effectiveness of RACT.

'I.  Designation of PACT                                         •                 '

     The-Draft CTG states that a 98% weight reduction of VOC emissions from
     continuous vents is representative of RACT for Polymers and Resins
     Manufacture.  This level of emissions•reduction is more typical of LAER
     tor process emissions sources.  For example, the draft CTG is more strin-
     gent than the preliminary draft NS?S for SOCMI Distillation Units.  It is
     also more stringent than the 97% reduction in benzene emissions specified
     xn the Agency's proposed NSSEAP for Benzene from Maleic Anhydride"Manufacture.
     RACT retrofit Requirements for existing sources should be less stringent than
     the NSPS requirements for the same source category.  The RACT requirements  '
     should also be less stringent than BACT, LAER, and NESHAP requirements for
     similar Rypes of emissions.  The level of control specified by this CTG should
     be consistent with the levels specified in other CTG'-s that are under develop-
     ment for the control of VOC emissions.

     CMA recommends that RACT be' set at a lower percentage level of VOC emissions Deduc-
     tions that will allow individual states to select the level of emissions reductions
     for existing sources.  This permits the optimal selection of RACT used by the
     states to bring individual ozone non-attainment areas into compliance with
     the NAAQS.

II.  Flares as Equivalent RACT

     At the NAPCTAC public hearing on the Preliminary Draft CTG for Polymers and
     Resins Manufacture, CMA commented extensively on this issue.  Our concerns
     have not been addressed by the Agency in the latest draft of the CTG.  CMA
     maintains that the available data on flare destruction efficiency demonstrate
     that these devices qualify as RACT for VOC emissions control.  The Draft CTG
     should be changed to permit the use of flares as RACT for pblyolefins plants.

     CMA and EPA are currently funding a study of flare efficiency at the John Zink
     Company Test Center.  This study is designed to determine the VOC combustion
     and destruction efficiencies achieved by flares controlling -small continuous
     streams typical of those found in polymers and resins and other chemical
     •manufacturing processes.  Since EPA is intimately involved in this study,
     takes the position that any CTG provision that would preclude the use of
     flares is inappropriate and strongly recommends that the language in the
     final CTG not discourage their use.

     Furthermore, since flares are specified as an acceptable control technioue
     for meeting the NSPS for SOCMI Distillation Units, ,CMA sees no reason why
     flares should not be acceptable as RACT for the CTG for Polymers and Resins
     Manufacture.
CMA
                                       B-17
-------
  Mr. Jack R. Farmer
  June 21, 1982
  Page Three
III.  Model Plant Descriptions and Emissions Factors

      The model plant descriptions and emissions factors for polystyrene manufacture
      ar* generally not representative of current industrial practice.  A wide
      range of polystyrene polymers are manufactured, ranging from low molecular
      weight emulsions to high molecular weight crystalline polymers.  Emissions
      factors generally decrease with increasing molecular weight.            :
                                      1 I    • • : ,,            !'  Ml" 1,1    :      , ,,   j
      Hecent increases in plant steam costs have forced some resin manufacturers
      operating plants of capacities similar to the model plant to provide process
      vacuums using vscuur. pumps rather than steam eductors.  CXA provided similar
      ccrar.*r.ts tc EFA in a letter dated October 19, 1981, for the NSPS Polymers and
      Resir.s development activity.  In this letter we indicated that these process
      improvements have lowered typical emissions factors for continuous polystyrene
      units to 0.15 kg of VOC/1000 kg of resin or less.  This is approximately 5% of
      the model plant emissions factor of 3.09 kg of VOC/1000 kg of resin.  As a
      result, SPA has drastically overestimated total emissions from existing poly-
      sryrer.a facilities.  Using the more current emission factor, total industry
      wide process emissions are 11 MG of VOC/year rather than the 227 MG of VOC/
      year estimated by EPA.

      This reduction in emissions has significant impact on the cost effectiveness
      of the CTG.  At the lower emissions rate, more typical of industry practice,
      th« control costs resulting from incinerating these emissions are $6,586./MG
      of VOC destroyed rather than the $320./MG of VOC destroyed estimated by EPA.
      For this reason, CMA concludes that further control of these emissions as
      specified in the CTG is unjustified.  We recommend that continuous polystyrene
      facilities with emissions factors typical of those described by CMA be exempt
      from the CTG on the basis that they are already demonstrating RACT.

 IV.  Limitations in the Incinerator Cost Data

      Again, CMA commented extensively on this issue at the NAPCTAC meeting, in
      preparing the Draft CTG the Agency ignored our comments on the limitations
      of the cost data.  We wish to reemphasize that EPA sjhould use the more
      representative cost information in the Hydroscience data base in determining
      th* cost effectiveness of this CTG.                                     •

      We trust that these comments will be considered as you prepare the final version
  of the CTG.  Enclosed for your review, are copies of our previously submitted comments
  referred to in this letter.  If you have any questions or comments, please contact
  Janet S. Matey, Manager, Air Programs at (202) 887-1179.  We thank you for having
  the oppcrtunity to comment on the Draft CTG.
                                                                              i

                                         Sincerely yours,
  Enclosures
B-18
-------
                                             ATTACHMENT TO JUNE 27, 7982 LETTER
                                             FROM 6.V. COX, CMA TO J.R. FARMERSEPA.
                             STATEMENT            .    — -
                                OF
                       RICHARD A, SYMULSSKI
                           ON BEHALF OF
              THE CHEMICAL MANUFACTURERS ASSOCIATION

                            BEFOftE THE
                 NATIONAL AIR POLLUTION CONTROL
                 TECHNIQUES  ADVISORY COMMITTEE

                                ON
         PRELIMINARY DRAFT CONTROL TECHNIQUES GUIDELINE
                                FOR
               CONTROL OF VOLATILE ORGANIC COMPOUND
                             EMISSIONS
                             FROM THE
           MANUFACTURE OF HIGH DENSITY POLYETHYLENE,
                POLYPROPYLENE,  POLYSTYRENE RESINS
JUNE 2, 1981
                                    CHEMICAL MANUFACTURERS  ASSOCIATION
                                    2501 H STREET, NW
                                    WASHINGTON, D,C,  20037
                              B-19
-------
      MY NAME is RICHARD A,  SYMULESKI,  I AM THE ENVIRONMENTAL CONSERVAJII
 COORDINATOR FOR THE AMOCO CHEMICALS CORPORATION SUBSIDIARY OF STANDARD
 OIL OF  INDIANA,  I AM SPEAKING TO YOU TODAY ON BEHALF OF THE CHEMICAL
 MANUFACTURERS ASSOCIATION'S PROCESS EMISSIONS REGULATIONS TASK GROUP AND
                               I '          :• •    "'.' 'I  1   " ,!'  ;. ,   I   "r ,:; ,	.;l	IS if';.;;
 POLYMERS/RESINS WORK GROUP,  CMA is A NONPROFIT TRADE ASSOCIATION HAVING
 186 UNITED STATES MEMBER COMPANIES THAT REPRESENT OVER 90 PERCENT OF THE
 PRODUCTION CAPACITY OF BASIC INDUSTRIAL CHEMICALS WITHIN THIS COUNTRY,
 WE  ARE  PLEASED  TO HAVE THE  OPPORTUNITY TO PRESENT OUR VIEWS  ANp
 CONCERNS TO THIS COMMITTEE  ON THE PRELIMINARY DRAFT CONTROL TECHNIQUES
 GUIDELINE  FOR THE CONTROL OF VOLATILE ORGANIC COMPOUNDS FROM THE
 MANUFACTURE OF  HIGH DENSITY POLYETHYLENE, POLYPROPYLENE AND  POLYSTYRENE
 RESINS,  C7IA MEMBER COMPANIES HAVE A CONTINUING INTEREST IN  ENSURING •.
                              .-.',.';       ,..',.. M • f, ': . "l».  :"  \f •'• •  '•;' .' ; I   '•'•  "I.-'::''*:	:":
 THAT  EPA DEVELOPS,,  WHEN NEEDS WARRANT, CONTROL TECHNIQUES GUIDELINES
 (CTG'S) THAT ARE TECHNICALLY SOUND, REASONABLE, ADMINISTRATIVELY FEASIBLE|
                             *                   '        ' '      I
AND COST-EFFECTIVE.   CMA HAS ATTEMPTED TO WORK WITH THE AGENCY  IN
DEVELOPMENT OF  THIS CTG.  To DATE OUR EFFORT  HAS  BEEN LIMITED BY
EPA TO TRACKING THEIR REGULATORY  DEVELOPMENT  EFFORTS,   As A  RESULT,
WE HAVE COMMENCED A PROCESS TO THOROUGHLY REVIEW  THIS CTG AND HAVE
 INITIALLY  IDENTIFIED SIX MAJOR PROBLEM AREAS  IN THE PRELIMINARY
DRAFT, WHICH  WILL FORM THE  BASIS  OF MY PRESENTATION,   THESE  PROBLEMS
 INCLUDE;                    .                             •      !
        •  INACCURACIES IN THE ESTIMATION OF THE NUMBER OF PLANTS
           THAT COULD BE IMPACTED BY THE CTG,
        *  LIMITATIONS IN THE DESCRIPTION OF THE INDUSTRY, ITS
           EMISSIONS FACTORS. AND  IN THE DEFINED MODEL  PLANTS,
        *  LACK OF ADMINISTRATIVE FLEXIBILITY TO THE STATES IN
                                                               s
           IMPLEMENTATION OF REASONABLY AVAILABLE CONTROL
                                                               I
           TECHNOLOGY (RACT),
                               B-20
-------
                                 -2-
        • DEFINITION OF  INCINERATION  AS  RACT FOR  THESE  SOURCE
          CATESORIES AND EXCLUSION OF APPLICABLE  AND  APPROPRIATE
          ALTERNATE CONTROL TECHNIQUES,
        •  FAILURE TO  RECOGNIZE FLARES AS EQUIVALENT RACT,
        •  LIMITATIONS  IN THE  COST DATA USED To.jusTiFY-RACT,
        BECAUSE  THESE PROBLEMS ARE EQUALLY APPLICABLE  TO ALL THREE
 POLYMERS INCLUDED IN THE CTG, MY COMMENTS WILL PROVIDE  A GENERAL  OVERVIEW
 OF THE  NATURE OF  THE PROBLEMS IDENTIFIED,   To ILLUSTRATE SPECIFIC POINTS
 I WILL  PRESENT  EXAMPLES  FOR INDIVIDUAL  POLYMERS,   HOWEVER,  THESE  COMMENTS
 SHOULD  NOT BE CONSTRUED  AS APPLYING ONLY TO THE SPECIFIC POLYMERS
 MENTIONED SINCE THEY DO  CROSS PRODUCT LINES,
             !•  NUMBER OF PLANTS SUBJECT To THE CTfi.

        THE CTG  IDENTIFIES A TOTAL OF  17 PLANTS IN OZONE NONATTAINMENT
 AREAS THAT HAVE REQUESTED STATE IMPLEMENTATION PLAN (SIP) EXTENSIONS
 FROM THE 1982 COMPLIANCE DEADLINE.  THESE PLANTS  WILL,  THEREFORE,  BE
 AFFECTED BY  THE CTG,  CMA HAS SURVEYED  THOSE STATES WHICH HAVE  REQUESTED
 AN EXTENSION OF THE OZONE ATTAINMENT  DATE AND WE  HAVE IDENTIFIED  A
 MINIMUM OF 25 PLANTS THAT WILL BECOME SUBJECT TO  THE  CTG.   THERE  ARE
 AN ADDITIONAL 23  PLANTS  LOCATED IN OZONE NONATTAINMENT  AREAS  FOR  WHICH
 NO EXTENSIONS HAVE YET BEEN FILED,   IT  IS HIGHLY  UNLIKELY THAT  ALL THESE
 AREAS WILL ACHIEVE COMPLIANCE WITH THE OZONE STANDARD BY THE  DEADLINE
 DATE,   THEREFORE,  A TOTAL OF  48 PLANTS COULD ULTIMATELY BECOME  SUBJECT
 TO THIS  CTG,
        IN ADDITION., THERE ARE  SEVERAL ERRORS  IN THE LIST OF PLANTS
 IDENTIFIED BY EPA  AS BEING SUBJECT tO THE  CTG,  AMOCO HAS PUBLICLY
ANNOUNCED THAT  IT  WILL NOT REBUILD ITS NEW  CASTLE  POLYPROPYLENE PLANT,
                                B-21
-------
                                 - 3 -

 ALSO., ACCORDING TO CMA DATA, THE BASF POLYSTYRENE PLANT USES A
 SUSPENSION POLYMERIZATION PROCESS WHILE THE RICHARDSON PLASTICS GROUP
 PLANT MANUFACTURES STYRENE COPOLYMER RESINS,  As A RESULT, THESE THREE
 PLANTS ARE NOT SUBJECT TO THE CT6,   FINALLY, WE WERE UNABLE TO CONFIRM
 EITHER THE LOCATION OR OWNERSHIP OF THE POLYSTYRENE PLANT IN DEE* PARK,
 TEXAS,

            II,  DESCRIPTION OF THE INDUSTRY, ITS EMISSIONS AND
                 MODEL PLANTS	

        IN DEVELOPING THE MODEL PLANTS FOR PURPOSES OF THE CT6, EPA
 RELIED SOLELY UPON INFORMATION FROM A SUBSET OF PLANTS LOCATED IN
 OZONE NONATTAINMENT AREAS.  THIS APPROACH SUBSTANTIALLY REDUCES THE    -
 NUMBER OF PLANTS USED TO DEVELOP PROCESS EMISSIONS FACTORS,  THE
 DRAWBACK TO THIS APPROACH IS THAT,  TO BE REPRESENTATIVE OF THE INDUSTRY
 AS A WHOLE, THE SELECTED SUBSET OF PLANTS MUST MIRROR THE UNIVERSE OF
 PUNTS IN EXISTENCE.   IF THIS -IS NOT THE CASE, THEN THERE IS SIGNIFICANT
 POTENTIAL-FOR INACCURATE DESCRIPTION OF PROCESS AND EMISSIONS CONTROL
 TECHNOLOGY IN PLACE,  AS WELL AS FOR THE PRODUCT SPECIFIC EMISSIONS
 FACTORS,   FOR EXAMPLE,  EMISSIONS FACTORS FOR HIGH DENSITYPOLYETHYLENE
 ARE BASED UPON PHILLIPS TECHNOLOGY,   HOWEVER,  MANYFIRMS MANUFACTURING
 THE PRODUCT EMPLOY MODIFIED  PHILLIPS TECHNOLOGY QR HAVE DEVELOPED
 SIGNIFICANTLY DIFFERENT PROPRIETARY PROCESSES  OF^ THEIR OWN,   THERE IS "
 SIGNIFICANT DIVERSITY IN THE HIGH DENSITY POLYETHYLENE PROCESS  TYPES
 IN  PLACE  TO WARRANT DEVELOPMENT OF  EMISSIONS FACTORS OVER  A LARGER
DATA BASE.          '             '. ' .         "'  	" ?""-!   '""lil!l  "'   -V    '^'^	
        IN THE  CASE OF POLYSTYRENE MANUFACTURING,  EPA HAS AGGREGATED
EMISSIONS DATA FOR TWO DIFFERENT PROCESS  TYPES  (BATCH-AND  CONTINUOUS)
TO DEVELOP THE MODEL  PLANT,  THE RESULT  IS A MODEL PLANT THAT OVERSTATES
                                 B-22
-------
 EMISSIONS FROM CONTINUOUS POLYSTYRENE UNITS.  IN THE BATCH PROCESS
 FUSITIVE EMISSIONS FROM REACTOR LOADING OPERATIONS CONSTITUTE A MAJOR
 SOURCE OF EMISSIONS, BUT THESE EMISSIONS ARE RELEASED OVER A SHORT PERIOD
 OF TIME, TYPICALLY 1 TO 2 HOURS EVERY 21 HOURS,  CONTINUOUS PROCESSES,
 ON THE ©TH€R HAND, HAVE CLOSES FEED SYSTEMS AND DO NOT HAVE AN EMISSIONS
 RATE FOR REACTOR LOADING OPERATIONS AS GREAT AS THE BATCH PROCESSES,
 APPLICATION  OF THESE DATA FOR BATCH PLANTS TO CONTINUOUS UNITS HAS RESULTED
.IN AN OVERESTIMATE OF THE REACTOR FEED EMISSIONS FACTOR.  ALSO, SINCE
 THESE EMISSIONS ARE ASSOCIATED WITH THE TRANSFER OF FEEDSTOCKS FROM
 STORAGE FACILITIES, THEY ARI TRULY STORAGE AND FUGITIVE LOSS EMISSIONS AND '
 SHOULD NOT BE SUBJECT TO CONTROL UNDER A CIS FOR PROCESS EMISSIONS,
        THE EMISSIONS FACTORS FOR THE VACUUM SYSTEM VENTS FROM PQLYSTYREIYE
 MANUFACTURE  ARE ALSO QUESTIONABLE.  IT APPEARS FROM THE MAGNITUDE OF
 THE NUMBERS  THAT THESE FACTORS ARE BASED UPON VENTS FROM STEAM EDUCTOR-
 TYPI VACUUM  SYSTEMS.  fflGH FEEDSTOCK COSTS HAVE RESULTED IN MANY MAJOR
 PLANTS SWITCHING TO VACUUM PUMPS FOR GENERATING PROCESS VACUUM,  IN
 THESE PLANTS, OVERHEAD FROM THE DEVOLATILI2ER AND THE STYRENE RECOVERY
 SECTION ARE  CONDENSED FOR RECOVERY RESULTING IN LOWER EMISSIONS FACTORS
 FOR THI VACUUM SYSTEM,
        THERE rs ALSO ONI GENERAL PROBLEM WITH APPLYING THE MODEL PLANT
 CONCEPT TO THE POLYMERS/RISINS INDUSTRY,   HOST POLYMERS/RESINS PLANTS
 CAN MANUFACTURE SEVERAL PRODUCTS WHICH COVER-A BROAD RANGE OF MOLECULAR
 WEIGHTS.  PROCESS OPERATING CONDITIONS AND, THEREFORE, EMISSIONS CAN BE
 EXPECTED TO  VARY FROM ONE PRODUCT TO ANOTHER.  THE MOST APPROPRIATE
 EMISSIONS FACTORS WOULD CONSIST OF RANGES, SO THAT PRODUCT/PROCESS
 VARIATIONS COULD BE ACCOUNTED FOR,
                                 B-23
-------
                                 - 5 -

        IN SUMMARY,  THE AGENCY,  BY RELYING UPON AN INNACCURATE
 REPRESENTATION OF THE INDUSTRY, HAS DEVELOPED QUESTIONABLE MODEL PLANT
 CONFIGURATIONS FROM WHICH UNCONTROLLED EMISSIONS FACTORS WERE OBTAINED,
 AS A  RESULT,  THE EMISSIONS FROM THE INDUSTRY HAVE BEEN OVERSTATED AND
 THE ACTUAL EMISSIONS REDUCTIONS THAT WILL BE OBTAINED BY~IMPLEMENTING
 THE CT6 WILL'-MOST LIKELY BE MUCH SMALLER,

            III.   LACK OF FLEXIBILITY  FOR  THE STATES  IN IMPLEMENTING
                  RACT.	

        THE CLEAN AIR ACT REQUIRES  THAT STATE IMPLEMENTATION  PLANS  FOR.
 NONATTAINMENT AREAS  MUST INCLUDE RACT REQUIREMENTS  FOR  STATIONARY
 SOURCES,  EPA HAS  PERMITTED STATES TO DEFER THE  ADDITION  OF  RACT
 REGULATIONS UNTIL  AFTER  THE AGENCY HAS DEVELOPED CTG's  FOR INDIVIpUAL
 SOURCE CATEGORIES.   THE CTG'S  ARE  TO PROVIDE STATE  AND  LOCAL AIR
 POLLUTION CONTROL  AGENCIES WITH AN INFORMATION BASE FROM  WHICH THEY
 MAY DEVELOP SPECIFIC RACT REQUIREMENTS.  THE CTG FOR  POLYMERS/RESINS
 MANDATES COMBUSTION  OF THE EMISSIONS  AT THE 98PERCENT  LEVEL AND AR6yES
 THAT THERMAL  INCINERATION  IS THE ONLY FEASIBLE APPROACH TO EMISSIONS
 CONTROL.  THIS APPROACH  SEVERELY CONSTRAINS  THE  STATE'S ABILITY TQ
'SELECT THE MOST  COST-EFFECTIVE CpNTROL STRATEGIES FOR STATIONARY
 SOURCES UNDER THEIR  JURISDICTION.                               '.[,
        IT is CMA's POSITION THAT IF A CTG as NEEDED FOR A SPECIFIC
 SOURCE CATEGORY, IT  SHOULD BE DEVELOPED CONSISTENT WITH THE GOALS
 AND REQUIREMENTS OF  THE  CLEAN  AlR  ACT,  As  SUCH,  THE  CTG  MUST BE ABLE
 TO DESCRIBE FOR  THE  STATES THE APPROPRIATE  RACT  TECHNOLOGIES APPLICABLE
 UNDER A VARIETY  OF SPECIFIC CIRCUMSTANCES WITHOUT DICTATING  TO THE
 STATES HOW TO CONTROL SOURCES  UNDER  THEIR JURISDICTION.   THE CTG SHOULD
 PROVIDE AN ACCURATE  DESCRIPTION OF^THE INDUSTRY,  ITS  EMISSIONS AND THE'
-------
                                 - 6 -
 TECHNOLOGIES CURRENTLY IN PLACE TO CONTROL THESE EMISSIONS,  FROM.THIS
 BACKGROUND THE CTG SHOULD DESCRIBE THE RACT's APPROPRIATE FOR THE
 SOURCES AFFECTED,  OVERALL GUIDANCE SHOULD BE GIVEN TO THE STATES IN
 INTERPRETING HOW SOURCES MIGHT BE AFFECTED BY THE CTG, BUT THE STATES
 MUST HAVE ULTIMATE CONTROL OF HO®, AND IF, THEY WILL IMPLEMENT THE CTG
 TO SOLVE THEIR SPECIFIC PROBLEMS,    THE CURRENT CTG DOES NOT FULFILL
 THESE OBJECTIVES,
            IV'   BgF INITTQH £F_lMC!MERATTQN A
        As CHAPTER IV OF THE PRELIMINARY DRAFT CTG STATES, RACT FOR
 THESE SOURCES HAS BON DEFINED AS 98 PERCENT CONTROL OF PROCESS VOLATILE
 ORGANIC COMPOUND (VOC) EMISSIONS WITH INCINERATION SERVING AS
 THE MODEL CONTROL TECHNIQUE,  THIS DEFINED LEVEL OF CONTROL FOR RACT
 IS EQUIVALENT TO THE LEVEL OF EMISSIONS CAPTURE OR REDUCTION NORMALLY
 DEFINED FOR  LOWEST ACHIEVABLE EMISSION RATE (LAER).  INCINERATION
 IS NOT GENERALLY EMPLOYED BY  THE INDUSTRY FOR THE CONTROL OF CONTINUOUS
 PROCESS VOC  EMISSIONS IN THE  POLYMERS/RESINS INDUSTRY,   FLARES,  HOWEVER,
 ARE THE MOST WIDELY USES CONTROL TECHNIQUE FOR BOTH CONTINUOUS AND
 INTERMITTENT EMISSIONS IN THE POLYMERS/RESINS INDUSTRY,   ON THIS BASIS,
 THE SELECTION OF INCINERATION FOR THE RACT STRATEGY IS  INCONSISTENT
 WITH THE GENERALLY ACCEPTED DEFINITION OF RACT,
        EPA'S  ARGUMENT FOR  EXCLUDING BOILERS AS RACT IS  INCONSISTENT WITH
 ITS  ARGUMENT  FOR  SELECTING  INCINERATORS,   WE AGREE WITH  EPA THAT BOILERS
AND  OTHER ENCLOSED COMBUSTION DEVICES  (INCLUDING  INCINERATORS) ARE  NOT
APPROPRIATE CONTROL DEVICES FOR  INTERMITTENT STREAMS,  SAFE,  EFFICIENT
OPERATION OF THESE DEVICES REQUIRES THAT THEY BE SIZED TO HANDLE  CONTINUOUS
                               B-25
-------
                                                 "Si I::!"
                                                 •!!, 1	"
                                 - 7 *
                                                                      ': 'If ('if11!!	.1 m	(Si
STREAMS WITH RELATIVELY STABLE FLOW RATES,  FOR EXAMPLE,  IlfTERMITTENT
LARGE VOLUME STREAMS  SUCH AS EMERGENCY RELEASES FROM REACTOR BECOMES It IONS
IN POLYOLEFINS PLANTS SHOULD BE CONTROLLED WITH FLARES,   IN  SELECTING
RACT, HOWEVER, EPA JUSTIFIED INCINERATION BY EXCLUDING  INTERMITTENT
RELEASES PROM CONTROL,  ON THIS BASIS, EPA'S ARGUMENT FOR EXCLUDING
BOILERS FROM RACT, EVEN THOUGH THEY HAVE BEEN  USED  BY INDUSTRY FOR THE
CONTROL OF  CONTINUOUS EMISSIONS, CANNOT BE SUPPORTED.
       OTHER NQNCOMBUSTION TECHNIQUES SUCH AS  ABSORPTION, ADSORPTION
AND CONDENSATION MAY ALSO PROVIDE HIGH VOC CAPTURE  EFFICIENCIES AT
                                 i !''  ' ' < i< ,v •'  ..... ............ " ....... i,J' !!"i,, ill ....... .......... ^.'i'51;"'..'!"1*;1  1'."'! "' ,:  ', I' . . ',*     I
LOWER COSTS THAN INCINERATION.  FOR THESE SOURCES THE DEFINITION OF
                 ,• i                I •,  ,«  ••',''  ''i"11"' :!'' '.„ ,„' I!1!!1,!;! ..... ::i I'v ,„!" J'1,,  , '  , ',< Jil, • ' ..... • I "„!  ., -     I
RACT SHOULD BE  FLEXIBLE ENOUGH TO  PERMIT  I NbuSTRY TO' SELECT THE MOST
COST-EFFECTIVE  MEANS OF CONTROL,
                                 . '  . .T ';   ;;"  ......... ;:.* ......  ','    , .......    .. ,| ...... ,' I ..... ....... ......... ,.T."
            V.   BLARES AS gQUTVALgMT  RACT
                                           .     ' ......... :   :,    .   •".      i "   .,'   , '. '.'.
                            .
       THE  DOCUMENT  IS INTERNALLY  INCONSISTENT IN ITS
JUSTIFICATION TO EXCLUDE  FLARES AS ACCEPTABLE RACT TECHNOLOGY,  tHE
DOCUMENT'S  ENVIRONMENTAL  ANALYSIS  OF  RACT STATES 'SINCE  FLARES WERE
                                 .:  .        '  ;  ;;""!!:!'J  ;:; ! :••";'..:. :'v? " v"'1' ..... :;::,:,;: :::l:T ......... •''
CONSIDERED AN ACCEPTABLE  CONTROL METHOD ONLY FOR INTERMITTENT STREAMS,
FLARES HAVE NO ROLE  IN THESE RACT RECOMMENDATIONS ,*  HOWEVER,  IN THE
DESCRIPTION OF EMISSION  CONTROL TECHNIQUES  THE DOCUMENT  STATES  "ELEVATED
FLARES HAVE A "ing CAPACTTV RANBP  ANTI ARP CAPABLE OF ADAPTING TO  CHANGSS
P
 IN EFFLUEflT FLQW  RATES  AND CONCENTRATIONS THAT ARE FOUND  IN  THE POLYMER
 INDUSTRY,"  CHA CONTENDS THAT A PROPERLY DESIGNED FLARE SYSTEM CAN
 HANDLE BOTH INTERMITTENT AND SMALL CONTINUOUS STREAMS, AND  PART OF THIS
 DOCUMENT AGREES WITH OUR CONTENTION,  IN FACT, FLARES ARE THE ONLY
 CONTROL TECHNOLOGY DESCRIBED IN THIS DOCUMENT WHICH  CAN HANDLE BOTH TYPES
 OF STREAMS,
                                  B-26
-------
                                 - 8 -
       THERE ARE  A NUMBER  OF  ENGINEERING  PRACTICES CURRENTLY IN USE
WITHIN INDUSTRY TO DEAL WITH  FLARING  LOW  FLOW CONTINUOUS EMISSIONS,  ONE
SUCH SYSTEM INVOLVES THE USE  OF  STAGED  ELEVATED FLARE SYSTEMS WHERE A
SMALL DIAMETER FLARE IS OPERATED IN TANDEM WITH A LARGE DIAMETER FLARE,
THE SYSTEM IS DESIGNED SUCH THAT THE  SMALL FLARE TAKES THE CONTINUOUS
LOW FLOW RELEASES AND THE  LARGER FLARE  ACCEPTS EMERGENCY RELEASES,   A
SECOND SYSTEM INVOLVES THE USE OF A SEPARATE  CONVEYANCE LINE TO THE
FLARE TIP FOR CONTINUOUS LOW  VOLUME,  LOW  PRESSURE RELEASES,   A THIRD
SYSTEM., SOMETIMES USED IN  CONJUNCTION WITH EITHER OF  THE ABOVE SYSTEMS,
INVOLVES THE USE  OF CONTINUOUS FLARE  GAS  RECOVERY,   IN THE LATTER SYSTEM
A COMPRESSOR IS USED TO RECOVER  THE CONTINOUOUSLY GENERATED FLARE GAS
"BASE LOAD."  THE COMPRESSOR  is  SIZED TO  HANDLE THE "BASE LOAD* AND ANY'
EXCESS GAS IS FLARED.
       A DISADVANTAGE L2STSD  IN  THE CTG FOR FLARE SYSTEMS IS THE „
POSSIBILITY OF DUCT FIRgS  FROM MANIFOLDING VENT STREAMS.   OBVIOUSLY,
THIS IS A DISADVANTAGE COMMON TO ALL  THE  DESCRIBED CONTROL
TECHNOLOGIES, INCLUDING THE PREFERRED RACT ALTERNATIVE,  INCINERATION,
       FLARE SYSTEMS HAVE  BEEN RECOGNIZED THROUGHOUT  THE POLYMER INDUSTRY
AS SAFE, COST-EFFECTIVE CONTROL  TECHNOLOGIES  WHICH  CAN ACHIEVE OR
APPROACH THE SAME DEGREE OF VOC  DESTRUCTION AS  OTHER  INCINERATION
DEVICES,   THE MOST DEFINITIVE DATA AVAILABLE  ON  FLARING  EFFICIENCIES
ARE CONTAINED IN THE GERMAN FLARE STU5Y BY  SlEGEL,  THE  RESULTS  OF  THE
GERMAN FLARE STUDY REPRESENT A YEAR'S WORTH OF  TEST DATA ON  FLARES  WHICH
CONSISTED OF ROUGHLY 1,300 TEST  SAMPLES.  THE TESTS WERE  PERFORMS)'AT
42 DIFFERENT MASS RATES, 23 DIFFERENT FLARE GAS  DENSITIES  AND  114 STEAM/
GAS RATIOS,   CONVERSION EFFICIENCY WAS  FOUND  TO  BE  INDEPENDENT OF MASS
FLOW,  WIND SPEED OR GAS COMPOSITION FOR THE REFINERY  GAS  STUDIED,   OF
                                B-27
-------
                                     •l.ii "*:';.; ,<\'.'i.
THE 1,300 TESTS ONLY  FOUR WERE LESS  THAN 99 PERCENT EFFICIENT,  THE
LOWEST VALUE (96 PERCENT) WAS  OBTAINED BY QUENCHINGTHE FLAME UNDER
CONDITIONS OF 10 TIMES  NORMAL  STEAM  RATE.  IN ALL CASES THE MEASURED
                _;,       • ,v .; ;. - '  ' '	| i;  i  • ;  •• r . .v;	<•  i^ i;: ,,i .   :/    •••• 	 '•'"•VV'T-'f
EFFICIENCY WAS GREATER  THAN 95 PERCENT,
       BATTELLE MEMORIAL LABORATORIES HAS CONDUCTED A STUDY FOR EPA TO
DEMONSTRATE MEASURING TECHNIQUES FOR USE AT FLARE TOWERS,  THE STUDY
WAS CONDUCTED OVER A  THREE-DAY PERIOD USING A JOHNZlNC FACILITY  FLARING
PROPANE.  ALTHOUGH THE  TEST HAS LONG BEEN COMPLETED, THE BATTELLE STUDY
HAS NOT BEEN MADE  AVAILABLE (EVEN IN DRAFT FORM) FOR.PUBLIC REVIEW AND
                .fll' |  "'i1     I,' „!»•""„     ! 'v ,' '   'i'!1 ', „ *,: , '" 1	 ,,!,'1||i|!' i'1" ,,"", ' :"' , liii'- ! •' ',,'!, ilffi'  . I   * f! •' :;i','i|	''  ' ,  ! •! I  • »!i	• "•'' 'I 'V
COMMENT,  WE HAVE  LEARNED,  HOWEVER,  THAT (ALTHOUGH NOT A SPECIFIC
OBJECTIVE OF THE STUDY) DATA EXIST DEMONSTRATING THAT THE FLARE SYSTEM
WAS ABLE TO ACHIEVE A DESTRUCTION EFFICIENCY  OF GREATER THAN  95 PERCENTy
EVEN WITH A SMOKING FLARE,   CMA CONTENDS THAT THE CTG SHOULD  NOT  BE
ISSUED IN FINAL FORM  UNTIL THE RESULTS OF THIS STUDY CAN BE
EVALUATED.
       IN ORDER FOR EPA TO  BE  CONSISTENT WITH THI? .SPIRIT, IF  NOT THE
EXPRESS LANGUAGE OF THE RECENTLY ISSUED EXECUITVIE ORDER No, 1229]^
(FEBRUARY U, 1981),  EPA is UNDER AN AFFIRMATIVEownTO ALLOW THOSE
CONTROL OPTIONS THAT  DATA DEMONSTRATE WILL ACHIEVE THE ENVIRONMENTAL
OBJECTIVES OF THE  REGULATION,  BUTATA LOWER COST TO INDUSTRY,  IN THIS
REGARD, THE AGENCY SHOULD NOT  PRECLUDE SUCH TECHNICALLY SOUND AND  COST-
EFFECTIVE CONTROL  TECHNIQUES,  UNLESS THE AGENCY  ESTABLISHES AN
ADMINISTRATIVE RECORD THAT  CLEARLY DOCUMENTS THEiSg pOST-EFFECTIVE
CONTROL TECHNIQUES WILL OFFSET A SIGNIFICANT ENVIRONMENTAL BENEFIT THAT
COULD OTHERWISE RESULT,
B-28-
-------
                                - 10 -
«   LIMITATIONS TM THg
                                                CnsT
        CMA'S CONCERN IS THAT THE INCINERATOR DATA USED IN SUPPORT OF
 THE CIS ARE OUTDATED,  INACCURATE AND INCONSISTENT WITH THE INCINERATOR
 COST INFORMATION PREPARED FOR THE AIR OXIDATION CT6 AND NEW SOURCE
 PERFORMANCE (NSPS)  ACTIVITIES,   THE CITED REFERENCE FOR THE POLYMERS/
 RESINS  CIS INCINERATOR COST ESTIMATES IS A REPORT PREPARED FOR EPA BY
 SARD,  INC,  IN DECEMBER 1978.   THE COSTS IN THIS REPORT HAVE BEEN UPDATED
 TO  DECEMBER 1977 BY SARD,  INC,  BUT THE PRIMARY SOURCES FOR BOTH DESIGN
 AND COST INFORMATION IN THI SARD REPORT DATE BACK TO 1972,  IN ATTEMPTING
 TO  ACCOUNT FOR INFLATIONARY EFFECTS, EPA HAS USED INDUSTRY ACCEPTED
 INFLATION FACTORS TO UPDATE THE COSTS TO JUNE 1980,   THE AGENCY SHOULD ..
 RELY ON. MORE RECENT COST INFORMATION WHEN IT IS AVAILABLE RATHER THAN
 ATTEMPT TO  RELY UPON COST  DATA  ESCALATED OVER A NINE YEAR INTERVAL.
 THESE MORE  RECENT COST DATA WERE DEVELOPED FOR EPA BY HYDROSCIENCE
 AS  PART OF  THE AIR  OXIDATION  CTS/NSPS EFFORT.   THE INFORMATION I'M. THE
 SARD REPORT ALSO DOSS  NOT  INDICATE HOW MANY INCINERATOR COST DATA
 POINTS  WERE USED TO DEVELOP THE COST CURVES.   FROM THE SHAPE OF THE
 CURVES  IT COULD  BE  ASSUMED  THAT THIY WERE EXTRAPOLATED FROM  ONE OR
 TWO  POINTS BY THE USE  OF SCALING FACTORS,
        BOTH THE INCINERATOR DESIGN AND COST INFORMATION FROM THE 6ARD
 REPORT  ARE  INCONSISTENT WITH  THE INFORMATION GENERATED BY HYDROSCtENCS
 FOR  THE AIR OXIDATION  REGULATIONS.   1HIS 1 NCONSISTENCY EXISTS DESPITE*"
 THE  SIMILARITIES IN THE  APPLICATION OF INCINERATOR TECHNOLOGY FOR
 CONTROL OF VOC'S  FROM  POLYMERS/RESINS AND AIR  OXIDATION UWITS,   THE
 POLYMERS/RESINS  CTS USES AS A BASIS,, THERMAL INCINERATION  AT  1500° F  IN
AN INCINERATOR HAVING A 0.5 SECOND  RESIDENCE TIME,  THE AIR  OXIDATION
                                B-29
-------
                                -21-
                                 i     • •    „  _ •   \::
CT6 PROPOSED INCINERATING GASES OF  SIMILAR HEAT AND VOC CONTENT AT
1600° F IN AN INCINERATOR WITH  A 0,75 TO 1,0 SECOND RESIDENCE TIME.  IN
GENERAL, THE MORE  CONSERVATIVE  CRITERIA IN THE AIR OXIDATION CT6 WILL
ASSUME COMPLETE  COMBUSTION  OF ORGANICS AND WILL RESULT IN A MORE
EXPENSIVE  INCINERATOR  HAVING HIGHER SUPPLEMENTAL FUEL REQUIREMENTS.
PUT ANOTHER WAY, IF  THE SIZING  AND  OPERATING CRITERIA FOR THE INCINERATOR
FOR AIR OXIDATION  CHEMICALS ARE CORRECT, THEN THE DESIGN CRITERIA AND
                • i"      J	• "    ",i' I"	• • I " "i  .. ,s *,!/', , . . ' 'ii,'."  i:!!;' ; ' ,  /  L i ,. • ' . :  i ; •.'."  ," : •»,"
COST ESTIMATES FOR INCINERATION IN  THE POLYMERS/RESINSC76ARE TOO  LOW.
A SIMPLE COMPARISON  OF TOTAL ANNUALIZED COST ESTIMATES INDICATES THAT
THE SARD REPORT  WILL PREDICT INCINERATOR COSTS THAT ARE ON THE AVERAGE
25 PERCENT TO 35 PERCENT LOWER  THAN THE COSTS OF THOSE PREDICTED BY THE
HYDROSCIENCE DATA  BASE.  Wg BELIEVE THE HYDROSCIENCE COST DATA ARE  MORE"
                                 i '   ' :       '"' .-,. ' ;,'• ' •! r .: :, r '• i   • :.' •,   . •  '.'    ' :,. MI 	:. ,'i i,':'.
REPRESENTATIVE Of  INDUSTRY  EXPERIENCE AND SHOULD BE USED AS THE BASIS
                           ,    ; •••! '   '.        • ' .   ;•; :/'  ' • i •     "'  •   ~ <•  '; ;•	;,
FOR DETERMINING  THE  COST EFFECTIVENESS OF THIS, CT6,
      '!'"        '                  „, ! :       ',,,'.  .mi1 111"„ "II'i' In 'I  "  ' , •    i     ',',',' ii .',.
       IN  SUMMARY, CMA HAS  IDENTIFIED SEVERAL KEY DEFICIENCIES WITH
THE PRELIMINARY  DRAFT  CT6 WHICH ADVERSELY AFFECT THE UTILITY OF THE
              '   •  *,     ,  _,     ''[''' 	 ! ' 	•' 	'	•' 	-! ' :fl!li;,:!'!' '!  '"' '•	;l ''<':'• ',.•'' :; • !-'ii:,! :!;",v
CTG-TO THE STATES, AND INCORRECTLY  ESTIMATE THE BENEFITS THAT WILL
                                                                  i
ACCRUE FROM ITS  IMPLEMENTATION,  CMA BELIEVES THAT  A MORE REALISTIC
ASSESSMENT OF THE  INDUSTRY, ITS EMISSIONS AND THE CONTROL TECHNOLOGY
IN PLACE CAN BE  USED TO DEVELOP A- CT6 WHICH WILL DEFINE COST-EFFECTIVE
RACT STRATEGIES  THAT ARE WORKABLE  FOR INDUSTRY, WHILE MEETING THE
NEEDS OF THE STATES  IN BRINGING NONATTAINMENT AREAS  INTO COMPLIANCE.
       CMA IS GRATEFUL FOR  HAVING  HAD THEOPPORTUNITY TO ADDRESS THE
COMMITTEE  TODAY.   OUR  WORK GROUP WILL REMAIN AVAILABLE TO WORK WITH
EPA IN DEVELOPING  A  REASONABLE, COST-EFFECTIVE CTG  FOR THE POLYMERS/
RESINS INDUSTRY.   THIS CONCLUDES  MY FORMAL STATEMENT,  I WILL ATTEMPT
TO ANSWER  ANY QUESTIONS YOU MAY HAVE CONCERNING MY  PRESENTATION.
                                 B-3Q
-------
                                                     June 21, 1982
                                       P. O. Box 3766
                                     Houston. TX T7OO1
            Mr. 3ack R. Farmer, Chief
            Chemicals and Petroleum Branch
            Emission Standards and Engineering Division (MD13)          '                 .  .
            Environmental Protection Agency
            Research Triangle Park
            North Carolina  27711

            Dear .vir. Farmer:                                                   •

                 EPA published  the "Control of Volatile Organic Compound Emissions from the
            Manufacture of nigh-Density Polyethylene, Polypropylene, Polystyrene Resins" and
            requested comments.   Our  comments  below  are  directed  at the high-density
            polyethylene and polypropylene manufacturing sections:                    ,

                -   '  .  1. -   A RACT of 98 weight percent reduction in VOC emissions  ;
                            from continuous vent streams is based on biased data from    "',
                            unrelated sources.      !    ~"    '.   "            !	

                  .1. .PaSe 3'6» Paragraph two states"... . that 98 percent destruction
               .  efficiency  is sometimes  achievable . . .", page 4-1, paragraph four ~
                 states thermal and  catalytic incinerators or boilers and process heaters
               ,   ...can achieve 98 percent VOC destruction efficiency .. ."/and page 3-
            .:   .10 paragraph four states ". .  . 98 percent efficiency for incinerator was
                .assumed (for the model  plant) . .  .".  A 98  percent efficiency  can"   •
             .   -P.r.obably be achieved, but to  maintain  this efficiency on a continuous or
                 average basis has not been.supported by data in this document.    •

                     •-. The  supporting data in Appendix A is based on incineration tests
                 on  waste vent streams  from  an  oxidative  butadiene  unit, maleic
                 anhydride units, an  acrylonitrile unit,  acrylic acid units, and lab-scale
                 tests.  The  organic compounds and concentrations  in the  plant scale
                 tests  are  different  than those found in units subject  to  tnis guideline
                 document and  as stated on  page A-22  the VOC  reduction efiiciencv
                 acmeved is unique  for each VOC compound.  Again in  the lab-scale
                 tests, real world efficiencies  are not comparable as stated on page A-2<*
                 cue -to  excellent  mixing in laboratory equipment  resulting  in high
                 reported efficiencies.                     *•                 •      •

                . '   " The  achievement  and  maintenance of a 9S percent  destruction
                 efficiency are further questionable as  Petrq-Tex spent $2.5 million on
                 an incinerator  which achieved  seventy percent efficiency.  Only after
                 thousands  of dollars in  improvements, was a  greater'than 98  percent
                 efficiency achieved.  The Rohm and Haas tests were made only when
                 the  production  unit was "operating  smoothly  and  the  combustion
                 temperature was at a steady state".
A D-V'S'ON
                                                B-31
         GUL»" O'L
S HOU&TOM Ci«n»
909 FmttHttt Sl»tC*
-------
                                                                              'ii; Ml	\m'rm '"fin
     The use of data from the  CTG for  air oxidation  processes is not
adequate  as a  basis for  RACT in  this  guideline  document,  and  98
percent  control efficiency is  not  achievable  on a  continuous , and
average basis.
                                    ,  •    • ' ' : •*' ,  ''' "'!!'  r-k •' i   :; '.:  •'
     2.    Discussion throughout the document infers, only  incinerators
           and boilers can achieve RACT.
     There was only one  reference to the use of process heaters  and
one reference to the use of refrigerated condensers ais possible methods
to achieve RACT.   In the reference to condensers, the efficiency  was
questioned.  There was  no reference to the processing of vent  streams
by other  units operated at  the site or by neighboring units, nor  was
mere reference to  combinations of control techniques such as the use
of  refrigeratea  concensers  followed by  flare  comoustion.     The
reiteration  c: tr.erma.  incineratior. suggests a specific control method
anc tne autnority to specify control methods'is questionable.

      3.    The  documentation  of flare   efficiency  is not  complete,
           disregards  or   simplifies   efficiency  data,   and   makes
           inaccurate statements in flare  design anc operations.
                      •	  '     ..'.'• i/'i- , . "i1 ' .. ''..',   : *	i '•!'»'!'•!• .'^(•'•"•f1;. ": „ '	' 	 ' '""-'
      Page  **-! states flare efficiency ". .  .  cannot be  quantified in
absence  of  adequate test  data . .  .",  yet  four  flare  studies were
reviewed  with one study containing 1298 test measurements.  A portion
of this data was  disregarded  because the flared vent gas was from a
petroleum industry.  A petroleum industry vent gas Is'more'variable in
composition which makes  design more difficult and is more susceptible
to efficiency problems than a plastics industry flare.

      The description of a polymer plant flare on page 3-18 states, "The
flares are mainly used to  handle emergency biowdowns which  requires
the  control device  to  handle  large  volumes of  ga,ses  with  variable
coTiposiTions." This is true for high pressure processes, but not for low
pressure,   liquid  pnase  high-density  polyethylene and  polypropylene
manufacture.   On  page  3-12  good combustion design  for flares  is
questioned  aue  to  lack  of  "completely well-defined"  measurement
methods.    Agency or  society  approved methods give  reliability to
measurement methods, but the lack of these methods should not detract
from tne evaluation of combustion design or the merit  of the  flare
efficiency cata.  The phrase on  page 5-18 in reference  re flares stating.
". .  . variations  in  flow and  heat content of the waste strearn coulc
extinguish tne flame .  .  ." is completely false as flares'are designed
with continuous pilot flames.   A continuous pilb't flame  is essential to
ensure safe conditions.
         ,•   ,   •  '    •         ;  .  •.  • '•  '••"'••. ' '	iiil*	j"*l:	i:-"j.' >''V;.i • ,,-
      fc.     Cost  calculations  for  thermal incinerator  installation need
            corrections and cost justification of retrofit is incomplete.
                             i,l'.' -" .-.'.• '••',  '''• 1 I'll!!!1'111' 'TFVs'l '•'« '."- V ; i" ''I ',.  .•
      On page 5-5 the escalation index needs  updating, the operating
 labor cost of $11.10/hr (including overhead)  is incorrect and should be
 $19.iG/hr (including overhead and benefits), and the interest rate  of 10
 percent should  be updated  to  the  18-20 percentrange.  The  cost
 anaivses die r.o't  inciuae the  cost of  a. filter svsterr.  uostream of the
rfl, • !' '',,1,,'i'i • ,4 I1" "
                              B-32
-------
     incinerator  which is • necessafy  to remove  polymers  and  entrained
     liquids.    Page  5-11  and  5-1.3  did  not contain  operating  labor,
     maintenance  labor,  and  electricity  costs   for  the  reciprocating
     compressor nor maintenance labor costs for manifolding.

           In  order  to assess  the  reasonableness of  retrofit (page 5-18), a
     flare efficiency  of 90 percent was used without justification, and "the
    /-Incremental cost  effectiveness  on  page 5-19 considered""only the
   .  annuaiized cost of the incinerator in the  calculations when annualized
   :  cost of the incinerator plus the manifold, plus a filter system-,-plus the-
   ... compressor should have been considered.


     If you have any  questions, please acdress them  to Ms. 3. F. Dey at 713-754-
4709 or to M.  R. Vyvial at 713-420-4296.
                                         3. R, Strausser
3RS/tls. ,.
                                    B-33
-------
                 Polysar Incorporated
                 Resin* Division
                 Process Development & Engineering Group
                 29 Main Street. P.O. Box 116. Leomtnster. Mass. 01453

                 Telephone (617) 537-9901
                 telex 710 347-1924
                                          July-.19,  1982      	
to.  Jack R.  Farmer, Chief      .         '
   Chemicals  and Petroleum Branch'
   Emission Standards and Engineering Division                       :    : '   .
United  States Environmental. Protection  Agency            ;                :
Office  of Air Quality Planning and Standards
Research Triangle Park, NC  27711                          .      '.'  .     ^   ,

Dear 'Sir: ,'    .            '    „'',,'.'„•'             '    ' -    ''••'•-.'

This letter is written in response-to  the Draft Control Technique Guideline
 (CTG) dated May 19, 1982 en tit led .."Control of Volatile Organic Compound  .
•Emissions from'the Manufacture of  High. Density Polyethylene, Polypropylene
 and Polystyrene Resins".       • • .  ;7,:, ./:••  .      •       :   '  ... .;   ,.,  -  !.

 The contents of this "letter are  intended  to provide ""assistance to the-EPA
 in drafting revisions to the  existing  CTG.         '                 ••'••

 Company Background    .     '        :'   •   '         ,-            •

 Polysar is a'multinational manufacturer of rubber petrochemical  and thermo-
 plastic resins, with  annual  sales in 1981 of $1.3 billion.   Polysar currently
 operates three polystyrene manufacturing- facilities in the United States...
 Polysar1s'total production of polystyrene, plus our operating-experience  and
 technology at  these facilities,  enables Polysar to be a significant producer
 of North American  polystyrene.                                         .

 Principles of  Emission  Control for Polystyrene

 The  major raw  materials normally utilized in  the manufacture of polystyrene
 are  styrene, ethyl benzene (processing aid in process),  synthetic  rubber
 (impact  modifier)  and a high boiling plasticizer,  (mineral oil is  commonly
 used)..  "               .        .   .           *	--.,..     ,     .  ;    ,.

 Polystyrene  manufacture differs from the  production  of either polyethylene
 or  polypropylene in the type of equipment used.   One  of the prime reasons
 for these differences is the absence  of  highly volatile material (low boil-
 ing) in the  manufacturing process.                                      |

 Due to the  absenc6 of low boiling hydrocarbons in the system, condensation
 is  used as  a prime mode of recovering  unused  hydrocarbon from the vapor
 stream flashed off (or stripped)  from the product in  the last stage  of the

                                      B-34                                 '
-------
Mr. Jack R. Farmer.,  Chief
July  19, 1982
Page  - 2-
polymer manufacturing process.  The major components to condense are styrene
and  ethyl  benzene.   Their vapor pressures at 60eF and atmospheric pressure
are  3.5mms and S.Omms respectively, and this low vapor; pressure enables them
to be  readily condensed even at low concentrations in a vapor- stream. . Con-
densation  is believed to be an accepted mode of control in a  majority of the
bulk polystyrene facilities in use. in the United States.            "-  •

Monomer Recovery'           '      .          .          •      '          ..     ...

To achieve anv measure of profitability in the polystyrene industry,-it is ••'••'.  .
very important to. maximize the recovery of unused monomer.  Recovered monomer
can  normally be reused,, after purification, and must be successfully recycled
•to achieve profitable results.  :••;•'   ' •..   " ."..      '.,......' '  ..   ..--.'  "   •. , ..,..: .  " .;

Condensing unused monomer from the vapor phase which is'generated in the
last phase'of the process is,, therefore,'an absolute financial necessity..-   >.
to. the successful operation, of a polystyrene-process.       .      •':•'  -  :;   '  ;;
.Due to the physical properties of materials used.and the  type of:. process
 employed,  successive cool-ing/condensation steps  are used  to achieve, acceptable  "
 results both for proper recovery of the unused monomer and for environmental  .. .
.reasons."-- '.':. -^ •".• '•;'•.;../.-. v.;.'-.-.." '.:-.'•; ;"'-'£ -•;.''•'V-':-; .'-....'.;.- '•':-.'-'\-J":.:•"-''.  "'•.''-••• ^;-:v";."'-'AV'.. :.':^.•(';'• :'•.-.i-'

.Because "the main incentive for monomer recovery  is  financial,  it "is necessary' '••-.
 to. install adequate'condensation capacity to  achieve proper, monomer recovery.  ,,
 The .additional facilities required to achieve, good  environmental results by. ":  '.
 condensation, only .requires a small additional; capital expenditure.       .   •. .   :,'

The environmental.control at all our facilities  are based on condensation
 technology and satisfies the State agencies in Massachusetts, Ohio and North
 Carolina.   We are confident that condensation technology  will enable the
 emissions for polystvrene facilities to be reduced  to below the recommended
 RACT (Reasonably Available Control Technology) referred to in the CTG of
 0.3 kgs per 1000 kgs of product.    • ,'        "'      '                       .

 The attached information was taken from reports  which were approved -by the State
 •environmental agency in Worcester, Massachusetts, before  the construction of
 a new Polysar facility for polystyrene manufacture.  The  technology used in.
 this facility to achieve the emission  control levels shown is purely based    .
 on the condensation type technology discussed in this letter, and clearly
 illustrates the ability of this technology to achieve the control required.
 The data submitted is theoretical and will, be further substantiated during
 the initial operation: of the new  facility. "               /        :   .

 Based on the information submitted with this  letter concerning the emissions
 of the latest Polysar facility  in Leominster, Massachusetts, I believe it is
 readily apparent that the use of  an incinerator  is  not required due to the
 small volume of emissions actually discharged.

                                  -' 6-35
-------
Mr, Jack R. Farmer, Chief
July  19, 1982
Page -3-

In general, the installation and use  of  an  incinerator  is "an  excessive expense
in a polystyrene facility, both for purchase  and  installation of the original
equipment and for the  operating costs  of the  installed  equipment,  and this
is illustrated very clearly in the financial  analysis which is part of the
CTG.  The operation of an  incinerator for any polystyrene  facility would
incur a continuous purge of fuel gas  to  -maintain  the stability of  the flame.
This would  certainly incur a substantial use  of fuel gas on an annual basis
and would increase the-emissions of carbon  monoxide/dioxide and sulphur
compounds from the facility.   Due to  the nature of incinerators, their mode
of operation and the associated vapor collection  systems,  Polysar  believes.
that  an inadequately designed  and conceived installation could potentially
be hazardous to operate and would certainly have  significant  explosive risks
due to the  potential for collecting explosive hydrocarbons/air mixtures.
               .   •           \      •• V     •••'.'       . • •    .   .     •••]',•  ••'••
Summary     .      .       •, .    • ••  ,       '  • •    ' . •      • •     .    •.',;.:.'•;
                                         «           i              '     ''*'.•'
Due to the  nature  of the raw materials used in the polystyrene manufacturing  .
process, the use of incinerators for  emission control  is not required.

Condensation is the control technology most widely used in the industry.
This  technology is used to recover un-reacted monomer  (which is. recycled
in  the process) and to control emissions to within accepted levels.
 The needless installation and: use of an incinerator will burn fuel gas and
 will increase the emissions of carbon dioxide, carbon monoxide and sulfur
 dioxide.   The use of an incinerator and the associated vapor collection
 system,may inherently, be a potential explosion hazard.      •          •      '

 If you have any queries concerning the contents of this letter, or you wish
 to have further discussions on this matter., please do not hesitate to contact
 me.         '           •           : 	  '          • '•' '' :i'"";'     •     , -   ':	

                                 	  Very truly yours,  '

                                          POLYSAR  INCORPORATED
                                          RESINS DIVISION    •   '
                                          ,       •  »i
                                          Frank  Jy witrano,  Manager
                                          Process  Development  and  Engineering
 FJM/bb
 Att achment

 PRE 503-2.4
                                      B-36
                                                    :,' '!f!!!l ,;"' li'Uiiiiii ,„'' I '"'I''1'
-------
 Mr.  Jack R.  Farmer,- Chief
 July 19, 1982
 ATTACHMENT
                          POLYSTYRENE MANUFACTURING
              EMISSION CONTROL DATA TAKEN FROM REPORTS SUBMITTED
         TO CENTRAL MASSACHUSETTS AIR POLLUTION CONTROL BOARD~IK  1981
        SOURCE        :'

 Styrene 'Storage

 Recycle Styrene,Storage

 Ethyl  Benzene Storage  •

 Catalyst  Make-Up    :;

 First  Reactor     ...

.Main Reactor  ...  •- '"  •" :

 Vacuum Vent • -';•'. '• -••• , • ':'

 Finishing Area: Vent
'  • • " ' ANNUAL .    :
  AVERAGE TEMPERATURE

         608F

         606F

        .6o°F   ; , .  •

At  Process  Temperature

At  Process  Temperature!

At  Process  Temperature

At  Process  Temperature

,At' Process  Temperature
     EMISSION RATE
 KGS/1000 KG OF PRODUCT

        0.0183

        0.0060

       . 0.0004   . .  .

       . 0.0025  V  .;   '

,      .  0.0596; ••  .   .  .'

        0.0396 .   .

        0.0109    .

   :     0.0155:

  ;     .0;.1328 KGS
  .    / Pkk 1000 KGS   '
        OF PRODUCT  . ••
                                     B-37
-------
   DATE-


SUBJECT



   FKOM:



     TO-
           UNITED STATES ENVIRONMENTAL PROTECTION AGENCY

April 13, 1983 Office of Air Quality Planning and Standards
               Research Triangle Park, North Carolina  27711

Telephone Conversation with Dr. Gerald Madden, E.I., duPont de Nemours
and Company* Inc. (Phone No.:(609) 299-1120) Regarding the Use of Catalytic
Incinerators for Existing Polymers and Resins Plants.

James Berry, Chief
Chemical Applications Section, CPB, ESED

Polymers & Resins CTG File


SlftttARY;

     Dr. Madden wanted to assure that catalytic incinerators are not
preempted from the Polymers & Resins CTG.  He feels that the capital cost
of catalytic Incineration is not prohibitive and welcomes an opportunity
to prove 1t with a comparative study if there was a case for him to bid on.
Dr. Madden has made several presentations to ESED (including one on March 12, 1981
regarding the capabilities of the DuPont Torvex Catalytic Reactor.

     I said we would review the CTG and see if we felt strongly against
catalytic Incineration for this application.  If so, we would give him
the opportunity to disprove us.  Otherwise, we will change the CTG to
allow catalytic Incineration.
                                                          ;,': I s)
                                                                                     1     ill
       1320*4 (*•». 3-7*)
                                           B-38
-------
       APPENDIX C
MAJOR ISSUES AND RESPONSES
           C-l
-------
                APPENDIX C.  MAJOR ISSUES AND  RESPONSES
     The major issues raised in the comment  letters  on  the  May 1982
draft of the CTG document are summarized  in  this  appendix,  as  well  as
EPA's responses to the comments.   (The  comment  letters  themselves are
included as Appendix B.)  The major issues which  are discussed (and
the corresponding section of this  appendix)  are:   the inclusion of
flares as RACT (C.I); the acceptability of•other  control  devices, such
as condensers, catalytic incinerators,  absorbers, adsorbers,  and
process heaters (C.2); the  stringency of  RACT and 98 percent  VOC
reduction (C.3); the basis  of the  cost  analysis (C.4);  and  the scope
of the CTG regarding the inclusion of both poiystyrene (c.5)  and the
high density polyethylene liquid phase  solution process (C.6).  Minor
corrections or updates regarding the chemical reaction mechanisms of
emissions or the status of  individual plants were rectified without
                                                1  •'•}'< • 11  • i  "  "'' i   ''?",.'
further comment by EPA.
                              !''   ' .•         !    '?'•'.'[,'•;!-;:   ,;»"':":::"
C.I  THE INCLUSION OF  FLARES AS  RACT
     Summary of Comments:
     Several commenters  (the Texas Chemical  Council - TtC,  the Chemical
                              •                  '•''  ;.JB"t' '"' i1 • 'i  •  '' ""•'!'' i' ' • 'i •"    ! .•
Manufacturer's Association  -  CMA,  and  Gulf  Oil  Chemicals Co.  - Gulf)
were of the opinion  that flares  should  be included in RACT as equivalent
control to thermal  incinerators.   TCC  felt that flares should be
included, especially in  light  of recent tests by Battelle  and John
Zink,  Co. for  EPA  (Howes  et.al.,  Chapter 4,  Ref. 9).  CMA  noted the   i
forthcoming joint  CMA/EPA  flare efficiency study (using the methods
developed by Battelle  and  John  Zink,  Co.) and suggested that  any
language  precluding  the  use of flares would  be inappropriate  especial'ly
since  they were  already  acceptable for the SOCMI Distillation  NSPS.
Gulf remarked  that the quantified flare efficiency  results of  four
studies were  disregarded.   Gulf also disagreed with  several statements
regarding  flares  in the draft CTG:  (1) that polymer plant flares  are
                                 C-2                           :
-------
generally for large volume, variable composition emergency blowdowns
(p. 3-18) - which Gulf stated is true for high pressure processes, but
not for low pressure liquid phase polypropylene (PP) and high density
polyethylene (HOPE) production; (2) that good combustion design was
questioned due to lack of Completely well-defined measurement
methods (p. 3-12) - Gulf felt that lack of measurement methods should
not detract from design evaluation or the merit of efficiency data;
and (3) that variations in flow and heat content of the waste stream
could extinquish the flame (p. 5-18) - Gulf was of the opinion that
this statement was completely false since continuous pilot flames are
used for safety.
     Response:
     On the basis of the now available results from the joint CMA/EPA
flare testing (McDaniel, et al., Chapter 4, Reference 11), flares have
been included as RACT capable of achieving 98 percent VOC destruction
under certain conditions.  This study is the first to use the sampling
and chemical analysis method developed by Battelle for EPA and is the
first to test efficiency at a variety of non-ideal conditions where
lower efficiencies had been predicted.  (All previous tests had used
easily combustible gases that do not tend to soot.)  Although 98 percent
VOC reduction efficiency has been demonstrated only for certain prescribed
conditions of gas velocity and heat content, existing flares are
considered acceptable for RACT in light of the high heat content
streams (other than the product finishing streams) generally emitted
by the PP and HOPE liquid phase processes.
     With regard to Gulf's comments on statements about flares in the
CTG, it is agreed that the statement on p. 3-18 that flares are primarily
used for large, emergency releases is true for the polymer industry in
general, but not for low pressure, PP and HOPE liquid phase process.
The statement on p. 3-12, however, is that the individual effects of
time, mixing, and temperature on combustion efficiency could not yet
be evaluated because measurement methods were not completely well
defined.  This statement was not intended to infer that information
regarding design or efficiency was not available,, but that its extent
and value was limited.  The ongoing CMA/EPA tests use a method that
was developed based on the previous studies, and these ongoing tests

                                C-3
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are improving and expanding the data base.  The remark on p. 5-18
stated that extinguishing the flame of a flareis"conceivable" - not
likely.  Although it is true that flares with continuous pilots will
relight momentarily, all existing flares do not necessarily have
continuous pilots or automatic relighting systems such as have come
into general use in recent years.
                                                                      i	
C.2  ACCEPTABILITY OF CONDENSERS, CATALYTIC  INCINERATORS, ABSORBERS,
     AND  PROCESS HEATERS AS RACT
     Summary  of Comments                                              '
     Various  commenters were  of the  opinion  that  control devices  other
than thermal  incinerators  should  be  included as RACT.   Monsanto,  CMA,
Gulf,  and Polysar  commented that  condensers  were  more  appropriate RACf
for  polystyrene manufacture than  thermal  incinerators  because no
highly volatile material  is present  inpolystyrenemanufacturing  so
that condensation  is less  expensive  and already  in use by  the industry.
DuPont wanted to  ensure that  catalytic incinerators were not preempted
because DuPont feels the cost is  competitive withthermal  incineration.
Gulf and fCC were concerned  that the repeated mention  of thermal
 incineration implied that only thermal incineration was accepted  as
 RACT and that the States would, therefore, not allow alternative
 control methods such as condensers,  process heaters of combinations
                               I         ,,,'„'  ,|| |!!!||!'''!,"' !J":' |! !'„ " i 'I '"        .1 '   • ! ' '   i '• 'h i,i , 'i'i' " ' 11
 (Gulf), or absorption or other recovery''techniques  (TCC).   (The inclusion
 of flares as an alternate control technique was also suggested; this
 issue was discussed separately in that last section.)
      Response:             '                 '   '  '	 ''       	'	'	"	'	'	"_
      While the May  1982 draft focused on  thermal  incinerators it was
 not intended to give the  impression that  other control techniques
 except flares, which were then disallowed,  would  not  be capable  of,
 and thus  acceptable for,  achieving  98 percent reduction.   For example
 the May  1982 draft  CTG  (on p. 4-1)  set an emission  reduction of  98 weight
 percent  VOC  for polypropylene and high-density polyethylene  plants and
 an  emission  limit of 0.3  kg  VOC/Mg  polystyrene produced and  stated  |
 that  "other  control techniques such as refrigerated condensation that
 can achieve  the  same degree  of control should be considered  equivalent
 and acceptable."  The  May 1982 draft  also stated specifically that "combustion
                                  C-4
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control devices, such as thermal and catalytic incinerators or boilers
and process heaters, can achieve 98 percent VOC destruction efficiency"
required for polypropylene and high density polyethylene liquid-phase
processes.  Therefore, the current CT6 has been revised so that alternate
control techniques are clearly defined as acceptable RACT if they
achieve the appropriate emission reductions or limits.
     'In addition, this final CTG document has been revised so that
alternative control techniques are discussed, cost estimates are
presented for not only thermal incinerators but also flares for PP and
HOPE, and condensers for PS.
C.3  STRINGENCY OF RACT
     Summary of Comments
     TCC, CMA, and Gulf questioned the ability of thermal incinerators
to achieve 98 percent VOC destruction on a continuous or average basis
under normal and realistic design and operating practice.  Monsanto
specifically questioned the extrapolation of its incinerator test data
from acrylonitrile to polymer production, while Gulf questioned the
applicability of test data from the CTG for air oxidation processes.
TCC agreed that 98 percent VOC reduction was achievable in all new,
well designed and well operated incinerators, but believed RACT should
be based upon "demonstrated levels in equipment that operates pretty
much as designed without elaborate post installation modifications to
fine-tune it to maximum (efficiency) levels."  TCC also felt that
thermal incinerator efficiency should be discounted to more realistic
levels since flare efficiency was discounted.  CMA was of the opinion
that 98 percent reduction was more appropriate for LAER than RACT and
was not consistent with other VOC emission limits under development.
     Response
     The questions regarding the stringency of RACT and the capability
of thermal incinerators to realistically achieve 98 percent VOC destruction
probably have become superfluous since flares have been accepted as
capable of achieving equivalent destruction and they have lower cost
so that flares are likely to be used to satisfy RACT, where needed.
     However, the state of the art supports that new incinerators can
achieve 98 percent reduction if properly designed and operated, as TCC
                                C-5
-------
has acknowledged after discussion with vendors.  Also, Petro-tex
increased the efficiency of controlling emissionsfrom its Oxo Butadiene
process from 70 percent to over 98 percent through relatively low cost
(in comparison to total capital cost) modifications that improved
mixing.  Although the air oxidation process emission test data are of
value in assessing incineration capabilities in general, it is true
that they do involve other chemicals  and processes, some of which,
however, would be expected to  be more, not less, difficult to control.)
Since the May 1982 draft was published, final  results  became available
for EPA emissions testing at a polypropylene facility.  The results  of
this study  program showed VOC  destruction  efficiencies of mixtures  of
gaseous, liquid,  and solid wastes  greater  than 99.7percent for
temperatures of  1,600°F and greater and  1.5  seconds residence  time.
Therefore,  98 percent  should be  readily  attainable for 0.75  seconds
residence  time  since  kinetic studies show  that residence  time  is
beyond  0.5 to 0.75  sec is  not  a  determining  factor of reduction  efficiency
 (see  p.  D36).
      In order  to ensure that RACT is readily achievable at a reasonable
cost,  and  to avoid  giving  a competitive disadvantage  to already well-
 controlled facilities (another concern of the TCC),  existing incinerators
 and flares will  be considered  to achieve RACT without the need for
 modification or replacement.   Also,  in order  to prevent a potential
 safety hazard (from combusting high-oxygen content streams) and a  _
 potentially unreasonable cost effectiveness,  RACT for product finishing
 and product storage operations was changed from 98 percent to 0.35  kg
 VOC/Mg product from the extruder  on  in the manufacturing process
 (e.g., pelletizing and product storage).
      Regarding CMA's  comment  that RACT is not consistent with the
 other VOC  regulations  under development:  RACT is not more stringent
 than the  NSPS for PP  liquid phase and HOPE  liquid phase  slurry.  The
 only difference  with  the SOCMI  Distillation NSPS, besides the  SOCMI
 Distillation NSPS's anticipation  of the joint CMA/EPA flare testing,
 results is the  use of a total resource  effectiveness (TRE)  index.   The
 SOCMI  Air Oxidation  CT6 also  employs a  TRE  index  and exempted"streams
 already controlled by a thermal  incinerator.   The Polymers  and  Resins
 CTG  allows States  to  decide.whether to  require testing and  subsequent

                                  C-6                                j
-------
modification or replacement, based on a case-by-case analysis of cost
effectiveness.
C.4  BASIS OF COST ANALYSIS
C.4.1  Origin of Costs (GARD vs Enviroscience)
     CMA believes that the incinerator cost data used in preparing the
Polymers and Resin CT6 are "outdated, inaccurate, and inconsistent
with the incinerator cost information prepared for the Air Oxidation
CTG and New Source Performance (NSPS) activities."  Further, they
state that these Polymers and Resins CTG cost data, which were obtained
from the GARD report,* have been escalated over a 9-year period (from
1972), that the GARD report does not indicate how many data points
were used to develop the incinerator cost curves, and that the GARD
incinerator design basis (1,500°F combustion temperature and 0.5 second
residence time) differs from the Air Oxidation CTG basis (1,600°F and
0.75 to 1.0 seconds).  Also, CMA notes that the GARD report predicts
annualized costs that are 25 to 35 percent 1ower than those prepared
from the Enviroscience data.  Finally, CMA believes "... the Hydroscience
(now Environscience) cost data** are more representative of industry
experience and should be used as the basis for determining the cost
effectiveness of this CTG."
     Response:
     Because flares, not incinerators, will likely be the control
technology employed to meet the CTG emission limits, the CMA comment
is effectively academic.  Nonetheless, we feel it necessary to respond
to certain statements CMA made concerning the quality of the GARD
data.
     First of all, we disagree with CMA that the Hydroscience
(Enviroscience) costs are "more representative" than the GARD data.
In actuality, the GARD incinerator costs generally compare well with
 *"Capital and Operating Costs of Selected Air Pollution Control Systems."
   R.B. Neverill, GARD, Inc., Niles, Illinois.  EPA Report 450/5-80-002.
   December 1978.
**"0rganic Chemical Manufacturing Volume 4: Combustion Devices."
   IT Environscience, Knoxville, Tennessee.  EPA Report-450/3-80-026.
   December 1980.
                                C-7
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the Enviroscience costs.  For example, based on the  combustion  chamber
volume, the GARD purchase cost for a thermal incinerator  to  control
one set of vent streams in the CT6 is approximately  $65,000.  The
Enviroscience cost for an incinerator of  the same  size  is approximately
$54,000.  (Both costs are in June 1980 dollars.)   Thus,  the  GARD cost
is 20 percent higher.  Most of this difference  isdue to  the fact  that
the GARD cost includes  a low-pressure fan,  while  the Enviroscience
cost just includes the  combustion chamber.   Even  so, the differences
are small when compared to  the nominal  accuracy of the  CTG and  NSPS
estimates ( 30 percent).  From this, one  can conclude that the  two
sets of costs are  essentially  equivalent.
     CMA makes additional statements  regarding  the accuracy of  the
GARD data.  Our  responses to  these  points follow::
     1.  Although some older  references  were  used in preparing the
          text for the  incinerator section of  the report, the costs  in
          GARD are not  "nine  years  old."   In fact,most  of the data
          were obtained from  a 1976 EPA report prepared  by an  incinerator
          vendor.*  Additional  data were taken from quotations for
           incinerators  installed at GARD's affilfated corporation,
           GATX  Terminals.   In any case, more than 20 data points  were
           used  to prepare the curves.  They were  not ".  . .  .  extrapolated
           from one or two points by the  use of scaling factors,"  as
           CMA alleges.   Moreover, the Enviroscience costs are  not that
           much newer than the GARD, since  the  former are in  December
                            „ , ,i       „ .   " •   ' i1 	:,'i',, i, I';,,.',,!,"'' •	i,1: j .' T'  '.' „• "•	j.!" ,;•!,' •!,!,  '   ]:,-,'ii • f!;1:""! fluJl V1™,
           1979 dollars, while" the latter reflect  December 1977 data.
           (Indeed, if  CMA has any more current information —-  1982
           costs,  for instance—we would  look forward to  seeing it.)
           CMA provides no documentation  for the  25  to  35 percent
           difference between the GARD  and  Enviroscience annualized
           costs.  Nonetheless,  these differences approximate the
           ±30 percent  accuracy  range for these estimates.  Further,
           given the wide variation  in  the factors for  the operating
           and maintenance  costs and  capital charges, these differences
2.
  *"Report of Fuel  Requirements, Capital  Cost, and Operating Expense for
   Catalytic and Thermal  Afterburners," CE-Air Preheater/Industnal Gas
   Institute, Stamford, Conn.  EPA Report 450/3-76-031, September 1976.
                                  C-8                                 !
-------
          represent excellent agreement for such "study" estimates.
          Because these cost factors vary so widely, it is more meaningful
          to compare the purchased or installed capital costs than the
          annualized.
     3.    The differences between the Polymers and Resins and Air
          Oxidation CTG design parameters are relatively small and
          impact the purchase costs only about 20 percent.  (See
          Appendix B of Air Oxidation CTG and Appendix E of this CTG.)
          Moreover, the Enviroscience report "Control  Device Evaluation
          for Thermal  Oxidizers" (December 1979) states that a VOC
          control efficiency of 98 percent or greater is achievable
          with a 1,500°F combustion temperature4 the basis for the
          GARD costs.   Thus, it is likely, although less certain than
          for 1600°F,  that the GARD incinerator can also meet the
          98 percent emission reduction alternative listed in the CTG.
     To summarize:  The GARD and Enviroscience thermal incinerator
costs can both meet the costing requirements of the CTG.  The differences
between the two sets of costs fall within the accuracy limits of the
CTG estimates.  Further, the GARD data are as current and as well-founded
as the Enviroscience costs.  Because of this, it makes little technical
difference which costs are used in the document.  In deference to CMA,
thermal  incinerator costs in this version of the CTG are based on
Enviroscience.
C.4.2.  Cost Effectiveness Calculations
     TCC noted that existing control levels for eight HOPE slurry and
solution process plants varied from - 696 percent (already meeting
RACT) to +92 percent of the uncontrolled emission rate for the model
plants.  TCC, therefore, questioned the validity of the cost effectiveness
analyses, especially if the 98 percent reduction were based on uncontrolled
model plant levels.  Similarly, Monsanto was concerned that the cost
effectiveness of RACT for polystyrene would be unreasonably high for
plants that were already well controlled.
     Response:
     TCC's concern about the cost effectiveness of existing plants
with varying degrees of control is unwarranted.  The 98 percent reduction
                                C-9
-------
would be applied to uncontrolled emission levels for a particular
plant.  The cost effectiveness is assessed in general by using the
uncontrolled and expected existing control levels as the lower and
upper bounds of the analysis.  However, the streams already controlled
by an existing flare or thermal  incinerator are ho longer  required to
be controlled further.  Therefore, already well controlled plants
would be rewarded, even though the calculated cost effectiveness from
the revised cost analysis based  on Enviroscience  is  less than  $550/Mg
for a thermal incinerator and  less than $160 for  a flare if the  same
unit were  used to  reduce from  90 percent  (selected arbitrarily to
approximate the upper  end of the range  of existing control levels  for
which additional control might be required)  to  98 percent  reduction.
      Monsanto's concern regarding the high cost-effectiveness  of
control  for already  well-controlled  polystyrene manufacturing  facilities
was  evidently based  on the  misunderstanding  that  incineration  would  be
 required regardless  of existing  control  levels.  On  the contrary,  RA^T
 for polystyrene is defined  as  an emission limit that can be  met  by any
 combination  of  existing  and additional  processes  and controls.  However,
 thermal  incinerators were used as a worst case cost analysis even
 after the acceptance of the 0.3 kg/Mg emission limit that was  based on
 the use of a condenser.  If a plant already has low emissions, cost
 effectiveness  would not increase unreasonably because less control  or
 no control would be required to meet the emission limit.
 C.4.3.   Miscellaneous
      Gulf pointed out several suggestions about details in the cost
 analysis:  the escalation  index  needs updating;  (2) operating labor
 (including overhead)  should be  $19.10 rather than $11.10; (3) the
 interest  rate of  10 percent should  be updated  to 18-20  percent,
 (4) inclusion of  a  filter  system upstream of the incinerator  to remove
 polymers  and entrained liquids,  (5)  inclusion  of operating labor,   :
 maintenance labor,  and electricity  costs for the reciprocating  compressor
 and maintenance labor costs for manifolding,  (6) completion of  justification
 of  the  90 percent efficiency  used for  existing flares  in  assessing
 reasonableness of retrofit, and (7)  consideration of  total  system
 cost, not only incinerator cost, for incremental  cost effectiveness.
                                  C-10
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     Response:
     The following responses are made with respect to Gulf's comments
on details of the cost analysis:  (1) the revised analysis uses a
different and corrected escalation factor; (2) operating labor (including
overhead) is now $18 per hour;  (3) the interest rate (before taxes)
remains at 10 percent because the analysis is in real, constant dollars
not considering inflation (even with inflation, 18 to 20 percent
interest would be too high at the time of writing this final CTG
document), (4) the cost analysis still does not include a filter
system because the tested polypropylene plant incinerates liquids and
solid atactic waste along with  the gases and achieves greater than
99.7 percent VOC reduction; (5) the revised cost analysis includes
operating labor and electricity costs for the entire system and the
incinerator combustion chamber  and includes maintenance costs for
manifolding (under source legs  and ducts, fans and stack) and compression
(fan under duct, fan, and stack); (6) the 90 percent existing flare
efficiency, which was used to represent a range of flare efficiencies
of existing units (about 70-99  percent), is still used for worst case
cost effectiveness calculation  purposes to approximate the greatest
existing control efficiency for a device, other than a flare or thermal
incinerator, that might have to be replaced or augmented.  (It is not
even certain that all existing  flares meet the conditions known to
achieve 98 percent VOC reduction according to results of the joint
CMA/EPA test program even though they will be considered to satisfy
RACT requirements); and (7) the incremental cost effectiveness of the
revised cost analysis is correctly based on total system cost.
C.5  SCOPE OF CTG:  POLYSTYRENE CONTINUOUS PROCESS
Summary of Comments
     Monsanto, CMA, and Polysar questioned the need for polystyrene
production to be covered by the CTG.  CMA and Monsanto noted in a
October 19, 1981, submittal by  CMA regarding the NSPS that current
typical emission factors were 0.119 to 0.15 kg/Mg, which is about
5 percent of the model plant level of 3.09 kg/Mg, because of economic
incentives.  Polysar was also confident that condensation could be
used to meet the 0.3 kg/Mg emission limit.  All commenters were concerned
                                C-ll
-------
about the implication that thermal incinerators would be the required
control technology in spite of the relatively high cost effectiveness
(especially if calculated based on the lower emission factors).
     Response
     As discussed in C.2, the May 1982 draft CTG suggests the use of  |
condensers to control the major polystyrene vent streams.  Since a
mass emission limit per production rate is used, the cost effectiveness
would not become unreasonable for already well controlled plants.
Therefore, in accordance with the industry data and the revised cost
analyses for control of polystyrene processes by condensation  (which
resulted in costs of emission reduction ranging from $-700/Mg  to
$950/Mg considering a range of current emissions of 0.20 kg/Mg to
3.09 kg/Mg in steam and in air, and vendor or Enviroscience  costs)  the
emission level for polystyrene was reduced to 0.12 kg/Mg.  Although
most polystyrene  plants may already be attaining  the RACT emission
level, and the consequent  emission reduction may  be small, polystyrene
will be  retained  in  the  CTG to  ensure uniform  control  in nonattainment
areas  across the  country  so that  at least a minimum control  level  is  ,
applied  and  no unfair  competitive advantage  results.
C.6 SCOPE  OF THE CTG:   HIGH  DENSITY  POLYETHYLENE, LIQUID  PHASE
     SOLUTION  PROCESS  AND OTHER PROCESSES NOT CURRENTLY INCLUDED     '
Summary  of  Comment
     TCC presumed that the HOPE,  liquid  phase solution process was not
covered  by  the  CTG since the  model  plant was based on  the  slurry
 (particle form)  process and recommended, therefore,  that the CTG be
 revised  to  clarify that only  the slurry process is covered.
 Response
      The HOPE,  liquid phase solution process has been examined for the
 NSPS since the May 1982 draft CTG.  The solution process was concluded
 to be different from the slurry process in terms of emissions and
 control.  Therefore, the HOPE, liquid phase solution process  is not
 included in the CTG.                                                 j
      Analyses of emissions and control have hot been conducted and
 control  techniques guidelines and RACT have not been established for
 high-density polyethylene liquid phase solution processes or  for other
                                  C-12
-------
processes (e.g., polypropylene and polyethylene gas phase processes)
with a relatively small number of existing plants.  However, EPA may
subsequently analyze and establish control techniques for any or all
of such other processes.  In the meantime, a State may choose to
conduct its own model plant or case-by-case analysis and establish  its
own guidelines.
                                C-13
-------
i,,iuiii:i,,' in,,	,n   „,   .jay; :ii|ii|
-------
       APPENDIX D
EMISSION SOURCE TEST DATA
         D-l
-------
                                                                   '"My:1" jiirs
                APPENDIX D:  EMISSION SOURCE TEST DATA
     The purpose of this appendix is to describe the test results of
flare and thermal incinerator volatile organic compounds (VOC) emissions
reduction capabilities.  Background data and detailed information
which support the emission levels and reduction capabilities are
included.
     Section D.I of this appendix presents  the VOC  emissions test data
including individual test descriptions for  control  of process  sources
by flaring.  Sections  D.2 and D.3 present the VOC emissions  test data
for control of  process  sources  by thermal incineration  and vapor
recovery  system, respectively.   Section  D.4 consists  of comparisons  of
various VOC test results and  a  discussion exploring and evaluating  the
similarities and differences  of these  results.
D.I   FLARE VOC  EMISSION TEST  DATA
      The  design and  operating conditions and results  of the  five
experimental studies  of flare combustion efficiency that have been  !
conducted were  summarized  in  Section 3.1.1.1.   This section  presents
more detailed  results of  the  first flare efficiency emissions test io
                                I            , .,  |"  'I h||    I "„ , | '  ,    !  , '      I, '    ' n |(	
 encompass a variety of "non ideal" conditions  that can be encountered
 in an industrial setting.   These results represent only the first
 phase of an extended study of which a final report should be available
 by mid-1983.                                                         !
      The aforementioned experimental study was performed during a   :
 three week period in June 1982 to determine the combustion efficiency
 for both air-  and steam-assisted flares under different operating
 conditions.  The study was sponsored by the U.S. Environmental Protection
 Agency and the Chemical Manufacturers Association  (CMA).  The test
 facility and flares were provided by the John Zink Company.   A total
 of 23 tests were conducted on  the steam-assisted flares and  11 tests
                                  D-2
-------
on air-assisted flares.  The values of the following parameters were
varied:  flow rate of flare gas, heating valve of flare gas, flow rate
of steam, and flow rate of air.  This section describes the control
device and the sampling and analytical technique used and test results
for the steam-assisted flare.
D. 1.1  Control Device.
     A John Zink standard STF-S-8 flare tip was used for the steam-assisted
flare test series.  This flare tip has an inside diameter of 0.22 m
(8 5/8 in.) and is 3.7 m (12 ft. 3.5 in.) long with the upper 2.2 m
(7 ft 3 in) constructed of stainless steel and the long 1.5 m (5 ft
0.5 in) constructed of carbon steel.  Crude propylene was used as the
flare gas.  The maximum capacity of the flare tip was approximately
24,200 kg/hr (53,300 Ib/hr) for crude propylene at 0.8 Mach exit
velocity.  Variations in heating valves of flare gas were obtained by
diluting the propylene with inert nitrogen.
D.I.2  Sampling and Analytical Techniques
     An extractive sampling system was used to collect the flare
emission samples and transport these samples to two mobile analytical
laboratories.  Figure D-l is a diagram of the sampling and analysts
system.  A specially designed 8.2 m (27 ft) long sampling probe was
suspended over the flare flame by support cables from a hydraulic
crane.
     Gaseous flare emission samples entered the sampling system via
the probe tip, passed through the particulate filter, and then were
carried to ground level.  The sampling system temperature was maintained
above 100°C (212°F) to prevent condensation of water vapor.  The flare
emission sample was divided into three possible paths.  A fraction of
the sample was passed through an EPA Reference Method 4 sampling train
to determine moisture content of the sample.  A second fraction was
directed through a moisture removal cold trap and thence, into a
sampling manifold in one of the mobile laboratories.  Sample gas in
this  manifold was analyzed by continuous monitors for 0 , CO, C0_, NO
                                                       £        L.    X
and THC on a dry sample basis.  A third sample was directed into a
sampling manifold in the other mobile laboratory.  Sample gas in this
manifold was analyzed for SO,, and hydrocarbon species on a wet basis.
                                D-3
-------
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     Data collection continued for each test for a target period of
20 minutes.  Ambient air concentrations of the compounds of interest
were measured in the test area before and after each test or series of
tests.
     Flare emission measurements of carbon monoxide (CO), carbon
dioxide (C09), oxygen (C09), oxides of nitrogen (NOV), total hydrocarbons
           *-             £                         X
(THC) and sulfur dioxide (SO,,) were measured by continuous analyzers
that responded to real time changes in concentrations.  Table D-l
presents a summary of the instrumentation used during the tests.
D.I.3  Test Results
     Twenty three tests were completed on the steam-assisted flare.
Table D-2 summarizes the results of these tests.  The results indicate
that the combustion efficiencies of the flare plume are greater than
98 percent under varying condition of flare gas flow rate, including
velocities as high as 18.2 m/s (60 fps) flare gas, heat content over
11.2 MJ/m3 (300 Btu/scf), and steam flow rate below 3.5 units of per
unit of flare gas.  The concentrations of NO  emissions which were
                                            A
also measured during the testing ranged from 0.5 to 8.16 ppm.
D.2  THERMAL INCINERATOR VOC EMISSION TEST DATA
     The results of six emission tests and one laboratory study were
reviewed to evaluate the performance of thermal incinerators under
various operating conditions in reducing VOC emissions from the different
process waste streams generated during the manufacture of polymers and
several synthetic organic chemicals.  The variable parameters under
which the incinerator tests were performed include combustion temperature
and residence time, type of VOC, type and quantity of supplemental
fuel, and feedstocks (solid, liquid, and gaseous waste streams).  The
test results, which are summarized in Table D-3, in combination with a
theoretical  analysis indicate that high VOC reduction efficiencies (by
weight) can be achieved by all  new incinerators.
     Three sets of test data are available.  These are emission tests
conducted on (1) incinerators at polymers and resins plants by EPA,
(2) incinerators for waste streams from air oxidation processes conducted
by EPA or the chemical  companies, and (3) laboratory unit data from
tests conducted by Union Carbide Company on incinerated streams containing
various pure organic compounds.  (No adequately documented data were
                                D-5
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                                                                                                                             !,,;,,'iii, .n'l,	51,1, iinniiiii
-------
found for tests of incinerators at polymers and  resins  plants  that
were conducted by the companies.)
     The EPA test studies represent the most  in-depth work  available.
These data show the combustion efficiencies for  full-scale  incinerators
on process vents at four chemical plants.  The tests measured  inlet
and outlet VOC, by compound, at different  incineration  temperatures.
The reports include complete test results, process  rates, and  descriptions
of the test method.  The four plants tested by the  EPA  are:
     1.  ARCO Polymers, Deer Park, Texas,  polypropylene unit,
     2.  Denka Chemicals, Houston, Texas,  maleic anhydride  unit,
     3.  Rohm and Haas, Deer Park, Texas,  acrylic acid  unit, and
     4.  Union Carbide, Taft, Louisiana, acrylic acid unit*
The data from ARCO Polymers include test results based  on three different
incinerator temperatures and three different waste  stream combinations.
The data from Rohm and Haas also include results for three  temperatures.
The data from Union Carbide include test results based  on two  different
incinerator temperatures.  In all tests, bags were  used for collecting
integrated samples and a gas chromatagraph with  flame ionization
detector (GC/FID) was used for obtaining an organic analysis.
D.2.1  Environmental  Protection Agency (EPA) Polymers Test  Data2
     EPA conducted emission tests at the incinerator at the ARCO
Polymers, Inc., LaPorte polypropylene plant in Deer Park, Texas (listed
as ARCO Chemical, Co., in LaPorte, texas,  in the 1982 Directory of
Chemical Producers3)  to assess emission levels and  VOC destruction
efficiency.
     The ARCO polypropylene facility has a nameplate capacity  of
181,000 Mg/yr (400 million lbs/yr).3  The  facility  produces polypropylene
resin by a liquid phase polymerization process.  The facility  includes
two "plants" (Monument I and Monument II) comprised of a total of six
process trains producing a variety of polypropylene resins.  Both
plants discharge their gaseous, liquid, and solid process wastes to
the same incinerator  system where they undergo thermal  destruction.
The wastes in the plants occur from:
     a)   processing  chemicals and dilution solvents for the catalyst,
                                D-9
-------
     b)   spent catalyst,
                            ,   , ,„             •'•  ,,,il'i!,!'   ',  •• . I,'  i •   '  :  , •:'    "'I
     c)   waste polymeric material  (by-product  atactic polymer),  and
     d)   nitrogen-swept propylene  from  the  final  stages (product
          resin purge columns) of the  process.
                                     ' ' •        ,.- •  iJ"   .'!   [  ;.' :  :  "f '  . • , I ' • !• ;
The feed rates of these wastes to the  incinerator  vary according  to
which trains are running and what startups are  occurring in  the two
plants.  Feed rate variations were  observed  during the two weeks  of
the incinerator test.
     The waste heat  boiler  associated  with the  incinerator provides a
major portion of the process steam  needed by the two polymer plants.
                            ,!''!.,       .•,,•,',..  ,",	 ,  i'if;,  . „  , . •!! :	  ii     |
Natural gas  is used  as  an auxiliary fuel  to  fire the incinerator.  If
necessary, fuel oil  can  also be used.   Under full production conditions,
the atactic  waste provides  approximately 50  percent of the energy
needed to produce the steam, and  natural  gas use is reduced.
     D.2.1.1 Control  Device.   The  incinerator and associated equipment
were designed by John Zink, Company.  The system was put into operation
on August 16, 1978.   The incinerator's two main purposes are  to destroy
organic waste from  the polymer processes (primary) and to provide heat
to generate  steam  (secondary).   Figure C-2 depicts a flow diagram of
the incinerator  and associated equipment.  Each inlet stream  has its
own nozzle  inside  the incinerator.   Combustion air is fed into the
incinerator  at  the  burner nozzles located approximately 4 feet beyond
the  incinerator entrance.   The combustion air  flow rate is  regulated
manually.  The  quench air enters  the incinerator within 3 feet of the
burner nozzles.   It is used to maintain  a constanC temperature and
 provide excess  combustion air.  The quench air flow  rate  is  automatically
 regulated  by an incinerator temperature  controller.
      During  normal  operation with  all waste streams  entering  the      ;
 incinerator, the natural gas is cut back and the  atactic  waste becomes
 the major fuel  source.  The purge  gas, which has  a  low  fuel  value  because
 it is 95 percent nitrogen, is fed  continuously to  the incinerator  for
 destruction of the  VOC since there  is no gas storagecapacity in  the
 system.  During an  upset of the incinerator this  stream is  sent  to a
                     :•       ,1     ,       •'" :	 ,, 	I	•„,! ,','!',  I •   •"',,', ,, '!•"•, I " , ,
 flare.  ARCO provided data to illustrate normal  operating parameters
 of the incinerator.  These are listed in Table D-4 and  represent the
                                  D-10
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                                                       ''"I;;,»•»;'i';i	•;»	r1;;"
averages for the month of August 1981.  The followingare considered
                               i    ;   .•'••. : ;: • • :. iliiiii "Si r: • i :  '.".; >  "„ •  ,  .  , I
design parameters:
     a)   heat input  =2.18 MJ/s (7.45 x 106 Btu/hr),
     b)   air supply   =15.1 standard mVs at  O'C(33,900scfm, at 60°F)
     c)

     d)
          firebox temperature  =980°C average andi,200°C maximum (1,800°F
          average and 2,200°F maximum),
          firebox residence time  =1.5 seconds, and
     e)   pressure  =19 kPa (78 in. H20).
     D.2.1.2  Sampling and Analytical Techniques.  A secondary purpose
of the ARCO incinerator test was to compare resultsof different
analytical methods for to the measurement of VOC emissions.  During
the testing phase of this program, three differen£ methods were used
for the collection and analysis of hydrocarbons.  These were:
     a)   EPA Method 25,                                              '
     b)   Proposed EPA Method 18  (both on-site and off-site  analyses
          performed), and
     c)   Byron  instruments Model  90  sample collection  system  and
          Model  401 hydrocarbon analyzer  samplingsystem  and instrument
          combination.
     To characterize  the  VOC  destruction  efficiency  across the thermal
 incinerator,  liquid/solid, and gas  phase sampling was  performed.   The
 sampling  locations  were:
      a)    Incinerator inlet - waste  gas  stream
                            - natural gas stream
                            - atactic waste stream
      b)    Waste heat  boiler outlet,  and
      c)    Scrubber stack outlet (volumetric flow rate).
      The  sampling system used for Method 25 consisted of a mini-impinger
 moisture  knockout, a condensate trap, flow control system, and a
 sample tank.  Both pre- and post-sampling leak tests were performed to
 ensure sample integrity.  In  the case of Method 18, samples were
 collected using a modification of EPA Method 110 for benzene.  This
 modification was necessary due to the high moisture content of the
 incinerator gases and the positive pressure of th'e emissions.  To
 ensure that a representative, integrated sample was collected using
 the modified Method  18, three validation tests for sample flow rate
 and sample volume into the Tedlar bag were performed.
                                  D-14
-------
     The principle underlying the Byron method  is the same as  EPA
Method 25.  However, rather than using a modified standard GC, the
Byron method uses a process analyzer.  This  instrument speciates C?
from higher hydrocarbons, but gives a single value for all nonmethane
hydrocarbons.  After separation, all carbonaceous material is  combusted
to C02 which is then converted to CH4 before being measured by an FID.
Thus, the variable response of the FID to different types of organics
is eliminated in the Byron 401 as it is in EPA  Method 25.
     The oxides of nitrogen (NOX) content of the flue gas was  determined
using the methodology specified in EPA Method 7.  A detailed description
of all these sampling and analytical techniques can be found in the
ARCO test report.
     The total  flue gas flow rate was determined two or three  times
daily using procedures described in EPA Method  2.  Based on this
method, the volumetric gas flow rate was determined by measuring the
cross-sectional area of the stack and the average velocity of  the flue
gas.  The area of the stack was determined by direct measurements.
     The work performed during this program incorporated a comprehensive
quality assurance/quality control (QA/QC) program as an integral part
of the overall  sampling and analytical  effort.  The major objective of
the QA/QC program was to provide data of known  quality with respect to
completeness, accuracy, precision, representativeness, and comparability.
     D.2.1.3  Test Results.  The VOC measurements were made by at
least four of five independent methods for each of eight different
combinations of incinerator temperature and waste streams.  Table D-5
summarizes the  results of measured destruction  efficiencies (DE's) for
each of these conditions.
     The results indicate that the values for the DE's by Method 25
are consistently lower and of poorer quality.  The poorer quality is
indicated by the imprecision reflected by the much larger standard
deviations for  this measurement method.   The accuracy and representa-
tiveness of these values obtained from Method 25 is, thus, questionable.
If Method 25 results are disregarded, the DE's for all  testing combinations
are found to be consistently above 99 percent.
                                D-15
-------
r
                 Table  D-5.   ARCO  POLYMERS  INCINERATOR  DESTRUCTION  EFFICIENCIES  FOR EACH  SET
                                                           OF CONDITIONS
Percent Destruction Efficiency3



Conditions
AW/N3/WG
2,000'F
WHG/HG
1,800'F
AM/MG/HG
1,600'F
HG/VS
1,800'F
NG/HG
1,600*F
AH/KG
2,000'F
AH/HS
1,800«F
AM/KG
1,600'F


Method 18 (on-s1te)
HCC
>99.99777 ± .00008

>99.9979 * .0004

>99.99721 ± .00009

99.8 * .1

>99.76 * .07

99.99674 ± .00007

>99.990 ± .004

>99.9975 ± .0001

Calculated

Byron
THCd
99.994 ± .002

99.996 ± .001

99.9961 ± .0003

99.9 ± .1

99.8 ± .10

99.9941 ± .0001

99.983 ± .007

99.994 ± .002

for Each Method
' .'Ml,
Byron
NHHCe
99.997 ± .002

99.998 ± .001

99.9957 ± .0002

99.6 ± .4

99.88 ± .04

99.99796 ± .00005

99.983 ± .007

99.995 ± .003h

,, 	 	
, 	 ;' 1"; | i

Method 25f
99.844 ± .006

99.8 ± .4

99.6 ± .2

76 ± 20

66 ± 10

96.32 ± .08

" 98 ± 3 ' "

99 ± 1


Method 18 (off -site)
Speciated
HC3

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i
99.88 ± .04
1



• 1 	 •••-"-
99.9979 ± .0001

ln 	 ^ j.,4, 	 *.!,.. ~ttt~t«~~, = inn 	 gC in Stack gas — — — —
                    enBr foflowinfth^fsig^l^the standard deviation (statistically expectei true value would fall between  the
                   reported value minus the standard deviation and the reported value plus the standard deviation).
                  (33,900 scfra) air supply, —
                  19 kPa (78 1n. H20 pressure).
                 cHe*sured using proposed EPA Method 18  (on-site) for hydrocarbons (HC) utilizing gas  chromatography (6C) with a flame
                  foliation detector (FID).  The values with "greater than" signs (>) indicate that the VOC was below the detectable
                  I1«1t and tht detection level was used to calculate the DE's.
                 dawsured using the Byron  Instruments Model 90 sample collection system and the Bryon Model 401 Hydrocarbon Analyzer
                  SWpHng system and Instrument combination (utilizing reduction to methane and FID)  in the total  hydrocarbon (THC)
                  mode.                                            '                        	•'   '      "    ' !
                 "Measured using the Byron  Models 90 and 401 combination (utilizing reduction to methane' and FID) in the  nonmethane
                  hydrocarbon node.                                                                                         j
                 fMtasured using EPA Method 25 for total gaseous nonmethane organics (TGNMO) utilizing GC-FID.  Data not  believed to
                  represent true values.
                 9Bejsured using proposed  EPA Method 18  (off-site) for individual hydrocarbon species  utilizing GC-FID.
                 h01ff1cult1es with analysis - Based on  most probable value.
                                                                    D-16
-------
D.2.2  Environmental Protection Agency (EPA) Air Oxidation Unit Test
       Data
     The EPA test study represents the most in-depth work available
for full-scale incinerators on air oxidation vents at three chemical
plants.  Data includes inlet/outlet tests on three large incinerators.
The tests measured inlet and outlet VOC concentrations by compound for
different incinerator temperatures.  The referenced test reports
include complete test results, process rates, and test method descriptions.
The three plants tested are Denka's maleic anhydride unit in Houston,
Texas, Rohm and Haas's acrylic acid unit in Deer Park, Texas, and
Union Carbide's acrylic acid unit in Taft, Louisiana.  The data from
Union Carbide include test results for two different incinerator
temperatures.  The data from Rohm and Haas include results for three
temperatures.  In all tests, bags were used for collecting integrated
samples and a 6C/FID was used for organic analysis.
     D.2.2.1  Denka Test Data.4  The Denka maleic anhydride facility
has a nameplate capacity of 23 Gg/yr (50 million Ibs/yr).  Maleic
anhydride is produced by vapor-phase catalytic oxidation of benzene.
The liquid effluent from the absorber, after undergoing recovery
operations, is about 40 weight percent aqueous solution of maleic
acid.  The absorber vent is directed to the incinerator.  The thermal
incinerator has a primary heat recovery system to generate process
steam and uses natural gas as supplemental fuel.  The plant was operating
at about 70 percent of capacity when the sampling was conducted.  The
plant personnel did not think that the lowered production rate would
seriously affect the validity or representativeness of the results.
     1.  Control Device.  The size of the incinerator combustion
chamber is 204 m2 (2,195 ft2).  There are three thermocouples used to
sense the flame temperature, and these are averaged to give the temperature
recorded in the control room.  A rough sketch of the combustion chamber
is provided in Figure D-3.
     2.   Sampling and Analytical Techniques.  Gas samples of total
hydrocarbons (THC)S benzene, methane, and ethane were obtained according
to the September 27, 1977, EPA draft benzene method.  Seventy-liter
aluminized Mylar^ bags were used to collect samples over periods of
                                D-17
-------
12 ft
          (Inlet)
                                               FLOW'
                                             SIDE VIEW
                                             23 ft-3 Jin
            There are Three Thermocouples Spaced Evenly Across the Top of the Firebox.
            The Width of the Firebox is 6ft-6 in.
                                                                  "... ".ifi.
                       Figure  D-3.  Incinerator Combustion Chamber
                                                                                            17ft -Sin
                                                                                   (Outlet)
                                                                                  !!!;>,;r .  "I
                                                        D-18
-------
two to three hours for each sample.  The insulated sample box and bag
were heated to approximately 66°C (150°F) using an electric drum
heater.  During Run 1-Inlet, the rheostat used to control the temperature
malfunctioned so the box was not heated for this run.  A stainless
steel probe was inserted into the single port at the  inlet and  connected
to the gas bag through a "tee."  The other leg of the "tee" went to
the total organic acid (TOA) train.  A TeflonR line connected the bag
and the  "tee."  A stainless steel probe was connected directly  to the
bag at the outlet.  The lines were kept as short as possible and not
heated.  The boxes were transported to the field lab  immediately upon
completion of sampling.  They were heated until the GC analyses were
completed.
     A Varian model 2440 gas chromatograph with a Carle gas sampling
valve, equipped with matched 2 cm3 loops, was used for the integrated
bag analysis.  The SP-1200/Bentone 34 GC column was operated at 80°C
(176°F).  The instrument has a switching circuit which allows a bypass
around the column through  a capillary tube for THC response.  The
response curve was measured daily for benzene (5, 10, and 50 ppm
standards) with the column and in the bypass  (THC) mode.  The THC mode
was also calibrated daily  with propane  (20, 100, and  2000 ppm standards).
The calibration plots  showed moderate nonlinearity.   For sample readings
that fell within  the range of the calibration standards, an  interpolated
response factor was used from a  smooth  curve  drawn through the  calibration
points.  For samples above or below the  standards, the response factor
of the nearest standard was assumed.  THC  readings used  peak  height
and column readings used area integration measured with  an electronic
"disc" integrator.
     Analysis for carbon monoxide was done on samples drawn  from  the
same integrated gas sample bag used for  the THC,  benzene, methane, and
ethane analyses.  Carbon monoxide analysis was done  following  the  GC
analyses using EPA  Reference Method 10  (Federal  Register, Vol.  39,
No.  47,  March 8,  1974).  A Beckman Model  215  NDIR  analyzer was  used  to
analyze  both the  inlet and outlet  samples.
     Duct  temperature  and  pressure values  were obtained  from the
existing inlet port.   A thermocouple  was  inserted  into the  gas  sample
                                 D-19
-------
                                                 liiil	P'ii,"
probe for the temperature while a water manometer was used for the
pressure readings.  These values were obtained at "the conclusion  of
the sampling period.
     Temperature, pressure, and velocity values were obtained for the
outlet stack.  Temperature values were obtained by  athermocouple
during the gas sampling.  Pressure  and velocity measurements were
taken according to  EPA  Reference Method  2  (Federal  Register, Vol. 42,
No. 160, August 18,  1977).  These values also were  obtained  at  the
conclusion of the sampling period.
     3.   Test Results  -  The  Denka  incinerator achieved greaterthan
98  percent reduction at 760°C (1400°F)  and 0.6 second  residence time.
These results suggest that 98 percent control is  achievable by properly
maintained and operated incinerators under operating conditions less
stringent  than 870°C (1600°F) and 0.75 second.  Table D-6 provides a  [
summary of these  test results.
      D.2.2.2  Rohm and Haas  Test Data5.  The Rohm and Haas plant in
Deer Park, Texas, produces  acrylic acid and ester.  The capacity of
this facility has been listed at 181 Gg/yr  (400 million Ibs/yr)  of    '
 acrylic monomers.  Acrylic esters are produced using propylene,  air,
 and alcohols, with acrylic acid produced as  an intermediate.  Acrylic
 acid is produced directly from propylene by  a vapor-phase catalytic
 air oxidation process.   The  reaction product is  purified  in subsequent
 refining operations.   Excess  alcohol is recoveredand  heavy end  by-products
 are incinerated.   This waste incinerator  is "designed'to burn off gas
 from the two absorbers.  In  addition,  all"process"vents (from  extractors,
 vent condensers, and tanks)  that might  be  a potential  source  of  gaseous
 emissions are collected  in a suction vent system"W normally  sent to
 the incinerator.   An organic liquid stream generated  in the process is
 also burned, thereby providing  part of the fuel  requirement.   The
                                                                       i
 remainder is provided  by natural  gas.
       1.   Control  Device -  Combustion air  is added to the incinerator
 in an  amount to  produce six  percent oxygen in the effluent.  Waste
 gases  are flared during maintenance shutdowns and severe process
 upsets.   The incinerator unit operates at relatively shorter residence
 times  (0.75-1.0 seconds) and higher combustion temperatures (650° -
 850°C) [1200°-1560°F] than most existing incinerators.
                               i             ',,.",„ i|i.  •.' !   I •    ;	
                                                1 •• :,'  "'.f i  . i • •  . I"*   „	,
                                  D-20
                                           ii;,;.,	iiii j'ii
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-------
     The total  installed capital cost of the incinerator was $4.7 million.
The estimated operating cost due to supplemental natural gas use is
$0.9 million per year.
     2.   Sampling and Analytical Techniques - Samples were taken
simultaneously at a time when propylene oxidations, separations, and
esterifications were operating  smoothly and the combustion temperature
was at a steady state.  Adequate time was allowed between the tests
conducted at different temperatures for the incinerator to achieve
steady state.  Bags were used to colle'ct integrated samples and a
GC/FID was used for organic analysis.
     3.   Test Results - VOC destruction efficiency was determined at'
three different temperatures and a residence  time of  1.0  second at
each temperature.  The test results are summarized in Table C-6.
Efficiency  is found to  increase with  temperature arid, except  for  774°t
 (1425°F), is above 98 percent.   Theoretical calculations  show that
 greater efficiency would be achieved  at 870°C (16000F) and  0.75  second
 than at the longer residence  times  but lower  temperatures represented!
 in these tests.                                      /  .
     D.2.2.3  Union  Carbide  Corporation (UCC) Test Data6.  The total
 capacity for the UCC  acrylates  facilities  is  about 90 Gg/yr (200
 million Ibs/yr)  of  acrolein,  acrylic acid, and esters.  Acrylic acid ]
 comprises 60 Gg/yr  (130 million Ibs/yr) of this total.  Ethyl acrylate
 capacity is 40 Gg/yr (90 million Ibs/yr).   Total  heavy ester capacities
 (such as 2-ethyl-hexyl acrylate) are 50 Gg/yr  (110 million Ibs/yr).
 UCC considers butyl  acrylate a heavy ester.
      The facility was originally built in 1969 arid utilized British
 Petroleum technology for acrylic acid production,,  In 1976 the plant
 was converted to a technology  obtained under license from Sohio.
      1.  Control Device - The  thermal  incinerator is one of  the  two
 major control devices used in  acrylic  acid and acrylate  ester manufacture,
 The UCC incinerator was installed in  1975 to destroy acrylic  acid and
 acrolein vapors.  This unit was constructed  by John  Zink Company for
 an installed cost of  $3 million and  incorporates a  heat  recovery unit
 to produce  process steam at  4.1 MPa  (600  psig).  The unit  operates  ai
 a relatively constant  feed input  and supplements  the varying flow  and
                                  D-22
-------
fuel  value of the streams fed to it with inversely varying amounts of
fuel  gas.  Energy consumption averages 15.5 MJ/s (52.8 million Btu/hr)
instead of the designed level of 10.5 to 14.9 MJ/s (36 to 51 million
Btu/hr).  The operating cost in 1976, excluding capital depreciation,
was $287,000.  The unit is run with nine percent excess oxygen instead
of the designed three to five percent excess oxygen.  The combustor is
designed to handle a maximum of four percent propane in the oxidation
feed.
     The materials of construction of a nonreturn block valve in the
4.1 MPa (600 psig) steam line from the boiler section require that the
incinerator be operated at 650°C (1200°F) instead of the designed
980°C (1800°F).  The residence time is three to four seconds.
     2.   Sampling and Analytical Techniques - The integrated gas
samples were obtained according to the September 27, 1977, EPA draft
benzene method.
     Each integrated gas sample was analyzed on a Varian Model 2400
gas chromatograph with FID, and a heated Carle gas sampling valve with
matched 2-cm3 sample loops.  A valved capillary bypass is used for
total hydrocarbon (THC) analyses and a 2 m long, 3.2 mm (1/8-in.)
outer diameter nickel column with PORAPAK^ P-S, 80-100 mesh packing is
used for component analyses.
     Peak area measurements were used for the individual component
analyses.  A Tandy TRS-80, 48K floppy disc computer interfaced via the
integrator pulse output of a Linear Instruments Model 252A recorder
acquired, stored, and analyzed the chromatograms.
     The integrated gas samples were analyzed for oxygen and carbon
dioxide by duplicate Fyrite readings.  Carbon monoxide concentrations
were obtained using a Beckman Model 215A nondispersive infrared  (IR)
analyzer using the integrated samples.  A three-point calibration
(1000, 3000, and 10,000 ppm CO standards) was used with a linear-log
curve fit.
     Stack traverses for outlet flowrate were made using EPA Methods 1
through 4 (midget impingers) and NOX was sampled at the outlet using
EPA Method 7.
                                D-23
-------
     3.   Test Results - VOC destruction efficiency was determined at  ;
two different temperatures.  Table D-6 provides a summary of these
test results.  Efficiency was found to increasewith temperature.  At
(800°C) 1475°F, the efficiency was well above 99 percent.  These tests
were, again, for residence times greater than 0.75 second.  However,
theoretical calculations show that even greater efficiency would be
achieved at. 870°C  (1600°F) and 0.75 second  than at: the  longer  residence
times but lower temperatures represented in these tests,
     All actual measurements were made as  parts per  million  (ppm)  of
propane with the other  units reported derived from the  equivalent
values.  The values were measured by digital  integration.
     The incinerator  combustion  temperature for  the  first  six  runs was
about  630°C  (1160°F).   Runs  7  through  9  were madeat an incinerator
temperature  of about  800°C (1475°F).   Only during Run 3 was  the acrolein
process  operating. The higher temperature caused most of the compounds
heavier  than propane  to drop below  the detection limit due to the wide
range  of attenuations used,  nearby  obscuring peaks,  and baseline noise
variations.   The detection limit ranges from about 10 parts per billion
 (ppb)  to 10 ppm, generally increasing during the chromatogram, and
 especially near large peaks.  Several  of the minor peaks were difficult
 to measure.  However, the compounds of interest, methane, ethane,
 ethylene, propane, propylene, acetaldehyde, acetone, acrolein, and
 acrylic acid, dominate the chromatograms.   Only aceticacid was never
                              -	      '   'i     ',.  ••'(•' "	II f  " PV- "!,"!:  - -  •'•*•','?• ,, " | • , ..  ;.•
 detected in any sample.
      The probable reason  for negative destruction efficiencies for
 several light components  is generation by  pyrolysisfrom  other components.
 For instance, the primary pyrolysis products of  acrolein  are  carbon
 monoxide and  ethylene.  Except  for methane and,  to  a much  lesser
 extent, ethane  and propane, the fuel  gas  cannot'contribute  hydrocarbons
 to the outlet samples.
       A  sample taken  from  the  inlet  line knockout trap  showed  6 mg/g of
 acetaldehyde, 25  mg/g  of  butenes,  and  100 mg/g  of  acetone when analyzed
 by  gas  chromatography/flame ionization  detection (GC/FID).
 D.2.3  Chemical  Company Air Oxidation Unit Test Data
       These data are  from tests performed by chemical companies on
  incinerators  at two  air oxidation  units:   the Pefro-Tex oxidative    !
                                  D-24
-------
butadiene unit at Houston, Texas, and the Monsanto acrylonitrile unit
at Alvin, Texas.  Tests at a third air oxidation unit, the Koppers
maleic anhydride unit at Bridgeville, Pennsylvania,7 were disregarded
as not accurate because of poor sampling technique.8
     D.2.3.1  Petro-Tex Test Data9.  The Petro-Tex Chemical Corporation
conducted emission testing at its butadiene production facility in
Houston, Texas, during 1977 and 1978.  This facility was the  "Oxo" air
oxidation butadiene process.  The emission tests were conducted during
a period when Petro-Tex was modifying the incinerator to improve
mixing and, thus, VOC destruction efficiency.
     1.   Control Device - The Petro-Tex incinerator for the  'Oxo1
butadiene process is designed to treat 48,000 scfm waste gas  containing
about 4000 ppm hydrocarbon and 7000 ppm carbon dioxide.  The  use of
the term hydrocarbon in this discussion indicates that besides VOC,  it
may include nonVOC such as methane.  The waste gas treated in this
system results from air used to oxidize butene to butadiene.  After
butadiene has been recovered from air oxidation waste gas  in  an oil
absorption system, the remaining gas is combined with other process
waste gas and fed to the  incinerator.  The combined waste  gas stream
enters the incinerator between seven vertical Coen duct burner assemblies.
The incinerator  design incorporates flue gas  recirculation and a waste
heat boiler.  The benefit achieved by recirculating flue gas  is to
incorporate the  ability to generate a constant 100,000 Ibs/hr of
750 psi  steam with variable waste gas flow.10  The waste gas  flow can
range from 10 percent  to  100 percent of the design production rate.
     The incinerator measures  72 feet by 20 feet by 8 feet, with an
average  firebox  cross-sectional  area of 111 square feet.   The installed
capital  cost was $2.5  million.
     The waste  gas stream contains essentially no  oxygen;  therefore,
significant combustion air must  be supplied.  This incinerator  is
fired with natural gas which supplies 84 percent of  the firing  energy.
The additional  required energy  is  supplied by the  hydrocarbon content
of the waste  gas stream.   Figure D-4 gives a  rough sketch  of  this
unit.
                                 D-25
-------
                                          :: "i
WASTE
  GAS
                                  Augmenting
                                 (Supplemental)
                                  Air Duct
                         Redrculaticn
                            Air duct
                                                 RECJRCULAT1CIN
                                                    AIR FAN
             Figure  D-4.   Petro-Tex oxo  unit  incinerator.
                                         D-26
-------
     2.   Sampling and Analytical  Techniques.  Integrated waste gas
samples were collected in bags.  The analysis was done on a Carle
analytical gas chromatograph having the following columns:
     1.   6-ft OPN/PORASIL.R (80/100).
     2.   40-ft 20 percent SEBACONITRILER on gas chrom. RA 42/60.
     3.   4-ft PORAPAKR N 80/100.
     4.   6-ft molecular sieve bx 80/100.
     Stack gas samples were collected in 30 to 50 cc syringes via a
tee on a long stainless steel  probe, which can be inserted into the
stack, at nine different locations.  They were then transferred to a
smaller 1 cm3 syringe via a small  glass coupling device sealed at both
ends with a rubber grommet.  The 1-cm3 samples were injected into a
Varian 1700 chromatograph for hydrocarbon analysis.  The chromatograph
has a 1/8-in. x 6-ft column packed with 5A molecular sieves and a
1/4-in. x 4-ft column packed with glass beads connected in series with
a bypass before and after the molecular sieve column, controlled by a
needle valve to split the sample.  The data are reported as ppm total
HC, ppm methane, and ppm non-methane hydrocarbons (NMHC).  The CO
content in the stack was determined by using a Kitagawa sampling
probe.  The Og content ln *ne stack was determined via a Teledyne
^/combustible analyzer.
     3.   Test Results.  Petro-Tex has been involved in a modification
plan for its  'Oxo' incinerator unit after startup.  The facility was
tested by the company after each major modification to determine the
impact of these changes on the VOC destruction efficiency.  The incinerator
showed improved performance after each modification and the destruction
efficiency increased from about 70 percent to above 99 percent.  Table
D-4 provides a summary of these test results.  The modifications made
in the incinerator are described below.
November 1977
     Test data prior to these changes showed the incinerator was not
destroying hydrocarbons as well as it should (VOC destruction efficiency
as low as 70 percent), so the following changes were made:
     1.  Moved the duct burner baffles from back of the burner to the
front;
                                D-27
-------
     2.  Installed spacers  to  create a continuous slot for supplemental
air to reduce the air flow  through the burner pods;
     3.  Installed plates upstream of the burners so that ductwork
matches burner dimensions;
     4.  Cut slots in recycle  duct to reduce exit velocities and
improve mixing with  Oxo  waste  gas;                                     '
     5.  Installed balancing dampers in augmenting  (supplemental)  air
plenums, top and  bottom;
     6.  Installed balancing dampers in three of th'e five sections  of
the recycle duct  transition; and
     7.  Cut opening in  the recirculation duct to  reduce the  outlet
velocities.                                                             :
March  1978                                                             j
     After the  November changes were made,  a field  test was made in
December 1977,  which revealed that  the  incinerator VOC destruction
efficiency increased from 70.3 percent  to 94.1 percent.  However, it
still  needed  improvement.  After  much discussion and  study  the following
changes were  made in March 1978:                                       ;
      1.  Took the recirculation fan out of  serviceand diverted the
           ..,      ,      ,   . „   ;  i  '!'.. i   •   ..-.Mi	;h' "i" •"'";  i1,;-;	r"/''•;!'.•• '	"•<•:;
excess forced draft air into  the  recirculation duct;
      2.  Sealed off the 14-cm (5-1/2-in.) wide  slots  adjacent to the
burner pods and removed the 1.3 cm  (1/2-in.)  spacers  which' were installed
                               ;           '   	,'.. vii!,1"  !,   :     "'i	-'-I,    '{	it. • •,:.
 in November 1977;
      3.   Installed  vertical baffles between the  b'ottom row of burner
           '•                "        '••       '   ' • i"-'1 :i!i,,;-4";ji":l- 4	1'!' *U' '  """ :'••':.." t  „  ;..
 pads to improve mixing;
      4.   Installed  perforated plates between the live  recirculation
 ducts for better waste  gas distribution in  the incinerator; and
           1 i  ]   '      "i      '" 'I "'',',      '     , ll"' i'l I'1' '' 4 II !Jll	 "''I' " I'11" III I ' »  "  •'!': .if .  •: ,1' , .' ' ' | ,   'llh'li
      5.   Cut seven  3-in. wide slots in the recycle duct for better
 secondary air distribution.
 July 1978                                                             '
      After the March 1978 changes, a survey in April  1978, showed  the
 Oxo incinerator  to  be  performing  very well  (VOC destruction  efficiency
 of 99.6 percent) but with a high  superheat temperature of 450°C  (850°F).
 So, in July 1978, some  stainless  steel shields  were  installed over the
 superheater elements  to help lower the superheat  temperature.  A
                                  D-28
-------
subsequent survey in September 1978, showed the incinerator to be
still  destroying 99.6 percent of the VOC and with a lower superheat
temperature of 400°C (750 j).
     This study pointed out that mixing is a critical factor in efficiency
and that incinerator adjustment after startup is the most feasible and
efficient means of improving mixing and, thus, the destruction efficiency.
     D.2.3.2  Monsanto Test Data.H  Acrylonitrile is produced by
feeding propylene, ammonia, and excess air through a fluidized, catalytic
bed reactor.  In the air oxidation process, acrylonitrile, acetonitrile,
hydrogen cyanide, carbon dioxide, carbon monoxide, water, and other
miscellaneous organic compounds are produced in the reactor.  The
columns in the recovery section separate water and crude acetonitrile
as liquids.  Propane, unreacted propylene, unreacted air components,
some unabsorbed organic products, and water are emitted as a vapor
from the absorber column overhead.  The crude acrylonitrile product is
further refined in the purification section to remove hydrogen cyanide
and the remaining hydrocarbon impurities.
     The organic waste streams from this process are incinerated in
the absorber vent thermal oxidizer at a temperature and residence time
sufficient to reduce stack emissions below the required levels.  The
incinerated streams include  (1) the absorber vent vapor (propane,
propylene, CO, unreacted air components, unabsorbed hydrocarbons), (2)
liquid waste acetonitrile (acetonitrile, hydrogen cyanide, acrylonitrile),
(3) liquid waste hydrogen cyanide, and (4) product column bottoms
purge (acrylonitrile, some organic heavies).  The two separate acrylonitrile
plants at Chocolate Bayou, Texas, employ identical thermal oxidizers.
     1.   Control Device - The Monsanto incinerator burns both liquid
and gaseous wastes from the  acrylonitrile unit and is termed the
absorber vent thermal oxidizer.  Two identical oxidizers are employed.
The primary purpose of the absorber vent thermal oxidizers is hydrocarbon
emission abatement.
     Each thermal oxidizer is a horizontal, cylindrical, saddle-supported,
end-fired unit consisting of a primary burner vestibule attached to
the main incinerator shell.  Each oxidizer measures 18 feet in diameter
by 36 feet in length.
                                D-29
-------


                                               
                                               -------
                                                              CO

                                                              >1
                                                              in
                                                              O)
                                                              c
                                                              O
                                                              10
                                                              O
                                                             •<->
                                                              (O
                                                              s_
                                                              0)
                                                             i.
                                                             OJ
                                                            UD
                                                             I
                                                             
-------
                                                                          i' '""nil, ' Eh., Cilinfil'f, ITU!
     The thermal oxidizer is provided with special burners and burner
guns.  Each burner is a combination fuel-waste liquid unit.  The
absorber vent stream is introduced separately into the  top of the
burner vestibule.  The flows of all waste streams are metered and      !
sufficient air  is added for complete combustion.  Supplemental natural
gas is used to  maintain the operating temperature required to combust
the organics and to maintain a stable flame  on the burners during
rainiraum gas usage.  Figure D-5 gives a  plan  view of  the incinerator.
     2.   Sampling and Analytical  Techniques,  the vapor feed streams
(absorber vent) to the thermal oxidizer and  the  effluent gas  stream
were sampled and analyzed using a  modified  analytical  reactor  recovery
run method.  The primary recovery  run methods  are  Sohio Analytical
Laboratory procedures.
     The modified method involved  passing  a measured amount  of  sample  i
gas  through three scrubber flasks  containing water and catching  the
scrubbed gas  in a gas  sampling bomb.   The  samples  were then  analyzed
with a  gas chromatograph and the weight percent  of the components was
                               i' ". ,  . .     ......   ,   , '  ", ., :|: ; ''I 1v  ,'„> :',,  ':  „ < ' ;  I
determined.
                                   , "             . '.!'"  •' h   i  ;  ',:»!' -i"'"   •.  '   1
     Figure D-6 shows  the  apparatus  and configuration used to sample
the  stack  gas.   It  consisted of  a  sampling line  from the sample valve!
to the  small  water-cooled  heat exchanger.   The exchanger was then
connected  to  a  250  ml  sample  bomb  used to  collect the unscrubbed
                    .....        r          •       ' f t ' : "'I- 1"   :  '   ;• "•   , .......  |
sample.  The  bomb was  then connected  to .a pair of 250 ml bubblers,
each with  165 ml  of water  in  it.   The scrubbers, in turn, were connecied
to another 250  ml  sample bomb used to collect the scrubbed gas sample
which  is  connected  to a portable compressor.  The compressor discharge
then was  connected  to a wet test meter  that vents to the atmosphere.
      After assembling the apparatus,  the compressor was turned on
drawing the gas from the stack and through  the system  at a rate of
 about  90 cm3/s ( 0.2 ft3/min).  Sample  gas  was drawn until at least
 0.28 m3 (10 ft3) passed through the scrubbers.  After  the 0.28 m
 (10 ft3)  was scrubbed, the compressor was shutdown  and  the unscrubbed:
bomb was analyzed for CH4, C2's,
                                        and C3Hg,  the  scrubbed  bomb  was
 analyzed for N£, air, 02, C02> and CO, and  the  bubbler  liquid was     !
 analyzed for acrylonitrile, acetonitrile, hydrogen  cyanide,  and  total
                                  D-32
-------
organic carbon.  The gaseous samples were analyzed by gas
chromatography.
     3.  Test Results.  The Monsanto Chemical Intermediate Company
conducted emissions testing at its Alvin (Chocolate Bayou), Texas,
acrylonitrile production facility during December 1977.  The VOC
destination efficiency reported was 99 percent.  (Residence time
information was not available and the temperature of the incinerator
is considered confidential information by Monsanto.)
D.2.4  Union Carbide Lab-Scale Test Data12
     Union Carbide test data show the combustion efficiencies achieved
on 15 organic compounds in a lab-scale incinerator operating between
430° and 830°C (800° and 1500°F) and 0.1 to 2 seconds residence time.
The incinerator consisted of a 130 cm, thin bore tube, in a bench-size
tube furnace.  Outlet analyzers were done by direct routing of the
incinerator outlet to a FID and GC.  All inlet gases were set at
1000 ppmv.
     In order to study the impact of incinerator variables on efficiency,
mixing must first be separated from the other parameters.  Mixing
cannot be measured and, thus, its impact on efficiency cannot be
readily separated when studying the impact of other variables.  The
Union Carbide lab work was chosen since its small size and careful
design best assured consistent and proper mixing.
     The results of this study are shown in Table D-7.  These results
show moderate increases in efficiency with temperature, residence
time, and type of compound.  The results also show the impact of flow
regime on efficiency.
     Flow regime is important in interpreting the Union Carbide lab
unit results.  These results are significant since the lab unit was
designed for optimum mixing and, thus, the results represent the upper
limit of incinerator efficiency.  As seen in Table D-7, the Union
Carbide results vary by flow regime.  Though some large-scale incinerators
may achieve good mixing and plug flow, the worst cases will likely
require flow patterns similar to complete backmixing.  Thus, the
results of complete backmixing would be relatively more comparable to
those obtained from large-scale units.
                                D-33
-------
              Table D-7.  DESTRUCTION  EFFICIENCY  UNDER STAYED CONDITIONS
                  BASED ON RESULTS OF  UNION  CARBIDE  LABORATORY TESTSa
                                                                                   ',(!.,':„ i II .";;i'":,f,::Ti'«
Destruction Efficiency of Compound in Percent
at Residence Time
0.75 second
Flow b
Regime
Two-stage
Backmixing


Complete
Backmixing


Plug Flow


Temperature
1300
1400
1500
1600
1300
1400
1500
1600
1300
1400
1500
1600
Ethyl
Aery late
99.9
99.9
99.9
99.9 .
98.9
99.7
99.9
99.9
99.9
99.9
99.9
99.9
Ethanol
94.6
99.6
99.9
99.9
86.8
96.8
99.0
99.7
99.9
99.9
99.9
99.9
;• • 1 , ,i'L IJ'I ;,;, [
Ethyl ene
92.6
99.3
". 99.9
99.9
84.4
95.6
98.7
99.6
99.5
99.9
99.9
99.9
Vinyl
Chloride
78.6
99.0
99.9
99.9
69.9
93.1
98.4
99.6
90.2
99.9
99.9
99.9
0.5/1.5 sec
Ethyl ene
87.2/97.6
98.6/99.8
99.9/99.9
99.9/99.9
78.2/91.5
;93.7/97.8
98.0/99.0
;99.4/99.8
97.3/99.9
99.9/99.9
^ 99. 9/99. 9
99.9/99.9
aThe results of the Union Carbide work are  presented  as  a  series  of equations.   These
 equations relate destruction efficiency  to temperature, residence time,  and flow
 regime for each of 15 compounds.  The efficiencies  in  this  table were calculated
 from these equations.
bThree flow regimes are presented:   two-stage backmixing,  complete backmixing,  and
 plug flow.  Two-stage backmixing is  considered  a reasonable approximation of actual
 field units, with complete  backmixing and  plug  flow representing the extremes.
                                             D-34
-------
D.3  VAPOR RECOVERY SYSTEM VOC EMISSION TEST DATA13
     On July 14, 1980, Mobil  Company collected samples of hydrocarbon
emissions from the exhaust vent of the Vapor Recovery/Knockdown System
at its Santa Ana, California polystyrene plant.  The samples were
taken using a MDA-808 Accuhaler^ pump while velocity was determined
using a Kurz^ Model 441 air velocity meter.  Samples were taken while
the plant was in normal operation.  One set of samples was taken while
a vacuum was drawn on dissolver tanks.  Another set of samples was
taken while a vacuum was drawn on the flash tank.  Both sets of samples
were analyzed for styrene and ethyl benzene by an independnet laboratory.
Computations for emission rates were made based on velocity, sample
volume and sample time.  The test results, submitted by the company,
indicate that 0.942 kg/day of ethyl benzene and 10.018 kg/day of styrene
are emitted from the exhaust vent of the vapor recovery/knockdown
system.  No more information was provided regarding the sampling and
analysis procedure used by Mobil or the laboratory.  It is assumed
that standard industrial practices were used, thus generating valid
estimates of emissions.  However, the data should not be used as a
significant basis for emission limitation.
D.4  DISCUSSION OF TEST RESULTS AND THE TECHNICAL BASIS OF THE POLYMERS
     AND RESINS VOC EMISSIONS REDUCTION REQUIREMENT
     This section discusses test results as well as available theoretical
data and findings on flare and incinerator efficiencies, and presents
the logic and the technical basis behind the choice of the selected
control level.
D.4.1  Discussion of Flare Emission Test Results
     The results of the five flare efficiency studies summarized in
Section 3.1.1.1  showed a 98 percent VOC destruction efficiency
except in a few tests with excessive stream, smoking, or sampling
problems.  The  results of the Joint CMA-EPA study, summarized in
Table D-2,. confirmed that 98 percent VOC destruction efficiency was
achievable for  all tests  (including when smoking occurred) except when
steam quenching occurred within the range of flare gas velocities and
heating values  tested.  Flare gas velocities for the tests  reported
to date go up to a high of 18.2 m/s (60 fps) and lower heating values
go as low as 11.2 MJ/m3   (300 Btu/scf).  Additional testing is currently
                               D-35
-------
 being undertaken to determine the effect of higher  velocities,  in
 particular, on destruction efficiencies.
 D.4.2  Discussion of Thermal  Incineration Test Results               |
      Both the theoretical  and experimental  data concerning combustion
 efficiency of thermal incinerators are discussedin this section.   A
 theoretical consideration of VOC combustion kinetics leads to the
 conclusion that at 870°C (1600°F) and 0.75 secondresidence time,
 mixing is the crucial design parameter.14  Published literature indicates
 that any VOC can be oxidized to carbon dioxide and water if held at
 sufficiently high temperatures in the presence of oxygen for a sufficient
 time.  However, the temperature at which a given level of VOC reduction
 is achieved  is  unique for each VOC compound.  Kinetic studies indicate
 that there are  two  rate-determining  (i.e., critically slow) steps in
 the  oxidation of a  compound.  The first  slow  stepof  the overall
 oxidation  reaction  is the initial reaction in which the original
 compound disappears.  The initial reaction of methane (CHzj.) has been
 determined to be slower than  that of any other  nonhalogenated organic
 compound.  Kinetic calculations  show that,  at  870°C (1600°F), 98
 percent  of the  original methane  will react in  0.3  seconds.   Therefore,
 any nonhalogenated VOC  will  undergo  an initial  reaction stepwithin
 this time.  After  the  initial  step,  extremely rapid free  radical
  reactions  occur until  each  carbon atom exists as carbon monoxide  (CO)
  immediately before oxidation is  complete.   The oxidation  of CO is the
'  second slow step.   Calculations showthat, at 87(5°C (16006F), 98
  percent of an original  concentration of CO will react in 0.05second.
  Therefore, 98 percent of any VOC'would be expected to undergo the   !
  initial  and final  slow reaction steps at 870°C(1600°F) in about  0.35
  second.   It is very unlikely that the intermediatefreeradical reactions
  would take  nearly as long as 0.4 seconds to convert 98 percent of the
  organic molecules to CO.  Therefore, from a theoretical viewpoint, any
  VOC should  undergo complete combustion  at 870°C (1600°F) in 0.75
  second.   The calculations on which  this conclusion is! baseci have taken
  into  account the  low mole fractions of  VOC and oxygen which would  be
  found in  the actual system.  They have  also  provided for the great
  decrease  in concentration per unit  volume due  to the elevated temperature.
                                  D-36
-------
However, the calculations assume perfect mixing of the offgas and
combustion air.  Mixing has been identified as the crucial design
parameter from a theoretical viewpoint.
     The test results both indicate an achievable control level of
98 percent at or below 870°C (1600°F) and illustrate the  importance of
mixing.  Union Carbide results on lab-scale incinerators  indicated a
minimum of 98.6 percent efficiency at 760°C (1400°F).  Since lab-scale
incinerators primarily differ from field units in their excellent
mixing, these results verify the theoretical calculations and suggest
that a full-size field unit can maintain similar efficiencies if
designed to provide good mixing.  The tests cited in Table D-6 are
documented as being conducted on full-scale incinerators  controlling
offgas from air oxidation process vents of a variety of types of
plants.  To focus on mixing, industrial units were selected where all
variables except mixing were held constant or accounted for in other
ways.  It was then assumed any changes in efficiency would be due to
changes in mixing.
     The case most directly showing the effect of mixing  is that of
Petro-Tex incinerator.  The Petro-Tex data show the efficiency changes
due to modifications on the incinerator at two times after startup.
These modifications (see Section D.2.3.1, 3. Test Results) increased
efficiency from 70 percent to over 99 percent, with no significant
change in temperature.
     A comparison of the Rohm and Haas test versus the Union Carbide
lab test, as presented in Table D-8, indirectly shows the effect of
mixing.  The UCC lab unit clearly outperforms the R&H unit.  The data
from both units are based on the same temperature, residence time, and
inlet stream conditions.  The more complete mixing of the lab unit is
judged the cause of the differing efficiencies.
     The six tests of in-place incinerators do not, of course, cover
every feedstock.  However, the theoretical discussion given above
indicates that any VOC compound should be sufficiently destroyed at
870°C (1600°F).  More critical  than the type of VOC is the VOC
concentration in the offgas.   This is true because the kinetics of
combustion are not first-order at low VOC concentrations.  The Petro-Tex
                               D-37
-------
     Table D-8.  COMPARISONS OF EMISSION TEST RESULTS  FOR UNION  CARBIDE
             LAB INCINERATOR AND ROHM & HAAS FIELD  INCINERATORS
Compound
Rohm and Haas Incinerator
 Inlet         Outlet
(Ibs/hr)      (Ibs/hr)
                                                Union  Carbide  Lab  Incinerator
                                                   Inlet         Outlet	
(Ibs/hr)
                                                                (Ibs/hr)
Propane
Propylene

Ethane
Ethvlene
IM V 1 1 J 1 \* 1 I V*
TOTAL
900
'l800b

10
30
2740
150
15'ob
;
375
190
865
'•i' •;'.; 	 >•*,•'•? iv 	 .:
71.4
142.9
	 ,„ :,,,, :h ,'' i,! ,, ,1; 	 i1'. ,, i •, ,
0.8
2.4
217.5
1 'i, 	 :»; fly, ,',;"
0.64
5.6

""'3.9"""
3.4
13.54 '
Overall VOC
Destruction
Efficiency:
          68.4%
        93.!
 aTable  shows  the destruction efficiency of the four listed compounds for  the
  Rohm & Haas  (R&H)  field and Union Carbide (UC) lab incinerators.  The  R&H
  results are  measured; the UC results are calculated.  Both sets of  results
  are based on 1425°F combustion temperature and one second residence time.
  In addition, the UC results are based on complete backmixing and  a  four-step
  combustion sequence consisting of propane to propylene to ethane  to ethylene
  to C02 and HgO.  These last two items are worst case assumptions.
 bAre not actual values.  Actual values are confidential.  Calculations  with
  actual values give similar results for overall VOC destruction  efficiency.
                                    D-38
                                                 ;„;„;,!	ijia!":• SMsmha •. I •'•iAji^i.i.'iiki •;<
-------
results are for a butadiene plant, and butadiene offgas tends to be
lean in VOC.  Therefore, the test results support the achievability of
98 percent VOC destruction efficiency by a field incinerator designed
to provide good mixing, even for streams with low VOC concentrations.
     The EPA tests at Union Carbide and Rohm and Haas were for residence
times greater than 0.75 second.  However, theoretical  Calculations
show that greater efficiency would be achieved at 870°C (1600°F) and
0.75 second than at the longer residence times but lower temperatures
represented in these two tests.  The data on which the achieveability
of the 98 percent VOC destruction efficiency is based is test data  for
similar control systems:  thermal incineration at various residence
times and temperatures.  If 98 percent VOC reduction can be achieved
at a lower temperature, then according to kinetic theory it can certainly
be achieved at 870°C (1600°F), other conditions being equal.
     A control efficiency of 98 percent VOC reduction, or 20 ppm by
compound, whichever is less stringent, has been considered to be the
acheivable control level for all new incinerators, considering available
technology, cost and energy use.*4  This is based on incinerator operation
at 870°C (1600°F) and on adjustment of the incinerator after start-up.
The 20 ppm (by compound) level was chosen after three different incinerator
outlet VOC concentrations, 10 ppm, 20 ppm, and 30 ppm, were analyzed.
In addition to the incinerator tests cited earlier in this Appendix,
data from over 200 tests by Los Angeles County (L.A.)  on various waste
gas incinerators were considered in choosing the 20 ppm level.  However,
the usefulness of the L.A. data was limited by three factors:  (1)  the
incinerators tested are small units designed over a decade ago; (2) the
units were designed, primarily, for use on coating operations; and
(3) the units were designed to meet a regulation requiring only 90  percent
VOC reduction.
     The 10 ppmv level was judged to be too stringent.  Two of the  six
non L.A. tests and 65 percent of the L.A. tests fail this criteria.
Consideration was given to the fact that many of the units tested were
below 870°C (1600°F) and did not have good mixing.  However, due to
the large percent that failed, it is judged that even with higher
temperatures and moderate adjustment, a large number of units would
still not meet the 10 ppmv level.
                               D-39
-------
     The 20 ppm level  was judged to be attainable.  All  of the non L.A,
and the majority of the L.A. units met this criteria,  there was
concern over the, large number of L.A. tests that failed, i.e. 43 percent.
However, two factors outweighed this concern.
     First, all of the non L.A. units met the criteria.  This is
significant since, though the L.A. units represent many tests, they
represent the same basic design.  They all are small units designed
over a decade ago to meet a rule for 90 percent reduction.  They are
for similar applications for the same geographic  region designed in
many cases by the same vendor.  Thus, though many failed, they  likely
did so due to common factors and do  not represent a  widespread  inability
to meet  20 ppm.
     Second, the  difference between  65 percent failing  10 ppmv  and
43 percent failing  20  ppm  is larger  than  a direct comparison  of the
percentages would reveal.   At  20  ppm,  not only did  fewer  units  fail,
but  those that  did  miss  the criteria did  so  by a  smaller  margin and   •
would  require  less  adjustment.  Dropping  the criteria from  10 ppm to
20 ppm drops  the failure rate  by  20 percent, but  is judged  to drop the
overall  time  and cost for adjustment by over 50  percent.
      The difference between the two levels is even  greater when the
 adjustment effort for the worst case is considered.  The crucial point
 is how close a 10 ppm level pushes actual field unit efficiencies to
 those of the lab unit.  Lab unit results for complete backmixing     !
 indicate that a 10 ppm level  would force field units to almost match
 lab unit mixing.  A less stringent 20 ppm level   increases the margin
 allowed for nonideal incinerator operation, especially 'for the worst
 cases.  Given that an exponential increase  may occur in costs  to     ;
 improve mixing enough for  field units to approach  lab  unit efficiencies,
 a drop  from 10 ppm to 20  ppm  may decrease costs  to  improve mixing in
 the worst case by  an  order of magnitude.
       The  30 ppm  level was  judged  too  lenient.  The only  data indicating
 such  a  low efficiency was  from L.A.   All  other data showed  20  ppm.   •
 The  non-L.A.  data  and lab data meet 20 ppm  and the Petro-tex experience
 showed  that moderate  adjustment  can increase efficiency.   In addition,
 the  L.A.  units  were  judged to have poor  mixing.   The mixing  deficiencies
                                 D-40

                                                                          fci	        I
-------
were large enough to mask the effect of increasing temperature.  Thus,
it is judged that 20 ppm could be reached with moderate adjustment and
that a 30 ppm level would represent a criteria not based on the best
available control technology cost, energy, and environmental  impact.
                                D-41
-------
D.5  REFERENCES FOR APPENDIX D
 1.  McDaniel, M.  Flare Efficiency Study, Volume I.  Engineering-Science.
     Austin, Texas.  Prepared for Chemical Manufacturers Association,
     Washington, D.C.  Draft 2, January 1983.

 2   Lee, K.W. et al., Polymers and Resins Volatile Organic Compound
     Emissions from Incineration:  Emission Test Report, ARCO Chemical
     Company, LaPorte Plant, Deer Park, Texas, Volume I Summary of
     Results.  U.S. Environmental Protection Agency, Research Triangle
     Park, North Carolina.  EMB Report No. 81-PMR-l.  March 1982.
                                                  "•	         •     i
 3.  SRI  International,  1982 Directory of  Chemical  Producers.

 4  Maxwell, W. and G.  Scheil.  Stationary  Source  Testing  of a Maleic
     Anhydride Plant at  the Denka Chemical Corporation, Houston,
     Texas.   U.S.  Environmental Protection Agency,,  Research Triangle
     Park,  North Carolina.  Contract  No.  68-02-32814, March 1978.
                                                                     , i
  5  Blackburn,  J.   Emission Control  Options for  the  Synthetic  Organic
     Chemicals Manufacturing  Industry, Trip  Report.  U.S.  Environmental
     Protection  Agency,  Research Triangle Park,  North Carolina.   EPA
     Contract No.
     68-02-2577,  November 1977.

  6   Scheil   G.   Emission Control  Options for the Synthetic Organic  [
      Chemicals Manufacturing  Industry,  Trip  Report.  U.S.  Environmental
      Protection Agency,  Research  Triangle Park,  North Carolina.
      Contract No.                                                     :
      68-02-2577, November 1977.

  7.  Letter from Lawrence, A., Koppers Company, Inc., to Goodwin, D.
      EPA.  January 17,  1979.

  8   Air Oxidation Processes in Synthetic Organic  Chemical Manufacturing
      Industry-Background Information for  Proposed  Standard Preliminary
      Draft EIS.  U.S. Environmental Protection Agency.  Research
      Triangle Park, North Carolina.  August 1981.  p. C-7 and C-8.

   9.  Letter  from Towe,  R., Petro-Tex Chemical Corporation, to Farmer, J.,
      EPA.  August 15, 1979.                                          ;

  10  Broz, L.D. and R.D. Pruessner.  Hydrocarbon  Emission  Reduction j
      Systems Utilized by Petro-Tex.  Presented at  83rd National Meeting
      of  AIChE,  9th  Petrochemical and Refining Exposition,  Houston,
      Texas,  March 1977.)

  11.  Letter  from  Weishaar, M., Monsanto  Chemical  Intermediates  Co., to
      Farmer, J.,  EPA,  November  8,  1979.

  12  Lee,  K.,  J.  Hansen and  D.  Macauley.  Thermal  Oxidation  Kinetics
      of Selected  Organic Compounds.   (Presented  at the 71st  Annual
      Meeting of the APCA,  Houston, Texas, June 1978.)

                                  D-42
-------
13.  Letter and attachments from Bowman, V.A. Jr., Mobil  Chemical
     Company, to Farmer, J.R., EPA.  September 9, 1980.  p. 13-16.
     Response to Section 114 letter on polystyrene manufacturing
     plants.

14.  Memorandum from Mascone, D.C., EPA.  June 11, 1980.   Incinerator
     efficiency.
                                D-43
-------
-------
APPENDIX E:  DETAILED DESIGN AND COST ESTIMATION PROCEDURES
                          E-l
-------
                                                                          I'«:'ii'VV 'i1:"'-:!'/!
     APPENDIX E:  DETAILED DESIGN AND COST ESTIMATION PROCEDURES
E.I  GENERAL                                                          ;
     This appendix consists of a more detailed presentation of the
bases, assumptions, and procedures used to estimate"equipment designs
and corresponding capital and operating costs forflares, thermal
incinerators, catalytic incinerators, shell-and-tube condensers, and
piping and ducting.  The  basis of design  and cost  estimates are  presented
In the following sections:  E.2, flares;  E.3, thermal  incinerators;
E.4, catalytic  incinerators; E.5 shell-and-tube  condensers; and  E.6,
piping and ducting.  The  installation cost factors used  in each  analysis
and the  annualized  cost  factors  used in  all  of the cost  analysis are  :
given in Tables 5-2 and  5-3,  respectively.
E.2   FLARE DESIGN  AND  COST ESTIMATION PROCEDURE                       <
      Flares  are open combustion  devices  that can be used to  effectively
and  inexpensively  reduce VOC  emissions.   The polypropylene and polyethylene
industries  commonly use flares to control large emergency releases arid
some high VOC streams.  Elevated flares were costed based upon 60 fps
exit velocity and a minimum of 300 Btu/scf.   Flare height and diameter,
which are the primary determinants of capital  cost, are dependent on'
 flare flow rate, heating value, and temperature.  Associated piping  and
 ducting  from the process sources to a header and  from a header  to the
 flare were conservatively designed  for costing  purposes.  Operating  :
 costs for utilities were based on industry  practice (1  fps purge  of  ,
 waste gas plus natural gas for continuous flow  flare"; 80 scfh natural
 gas per  pilot, number of pilots based on flare  tip diameter;  0.4  Ib  .
 steam/1b hydrocarbon  at  maximum smokeless rate).
      E.e.l   Flare  Design  Procedure.  Design of  flare  systems  for  the
 various combinations  of  waste streams was basedonstandard  flare
 design  equations  for  diameter and  height presented by IT  Enviroscience.1
 These equations were  simplified  to  functions  of the  following waste  'gas
                                 E-2
                                                            •y',.:,,ijtti:r:,	u ,•'&'<'*,'.'iv.'iaswi.	-i	i	i"	M
-------
characteristics; volumetric flow rate, lower heating value,  temperature,
and molecular weight.  The diameter expression is based on the equation
of flow rate with velocity times cross-sectional  area.  A minimum
commercially available diameter of 2 inches was assumed.  The height
correlation premise is design of a flare that will not generate a
lethal radiative heat level (1500 Btu/ft2 hr, including solar radiation2)
at the base of the flare (considering the effect of wind).  Heights in
5-foot multiples with a minimum of 30 ft. were used.3  Natural gas to
increase the heating value to 115 Btu/scf is considered necessary by
vendors to ensure combustion of streams containing no sulfur or toxic
materials.4  A minimum lower heating value of 300 Btu/scf has been
shown to help ensure a 98 percent efficiency for steam-assisted flares.
For flares with diameters of 24-inches or less, this natural gas was
assumed to be premixed with the waste gas and to exit out the stack.
For larger flares, a gas ring was assumed if large amounts of gas were
required because separate piping to a ring injecting natural gas into
the existing waste gas is more economical than increasing the flare
stack size for  large diameters.  The  flare height and diameter selection
procedure is detailed in Table E-l.
      Natural gas was assumed at a rate of 80 scfh per pilot flame
to ensure ignition and combustion.  The  number of pilots was based on
diameter according to available commercial equipment.5  Purge gas also
may be required to prevent air intrusion and flashback.  A purge velocity
requirement  of  1 fps was  assumed during  periods  of continuous flow for
'standard systems without seals.6
      Steam was  added to produce smokeless  combustion  through  a combined
mixing and quenching effect.  A steam ring at the flare tip was used
to  add steam at a  rate of  0.4 Ib steam/1b  of  hydrocarbons  (VOC plus
methane and  ethane)  in the continuous stream.''   Availability and
deliverability  of  this quantity of  steam was  assumed.
      Piping  (for flows less  than 700  scfm) or  ducting (for  flows equal
to  or greater than 700 scfm) was designed  from the  process  sources to
a header  combining the streams  and  from  the  header  to the base of  the
 flare.  Since  it  is  usual  industry  practice,  adequate pressure  (approximately
3 to  4  psig) was assumed  available  to transport  all waste gas  streams
without use  of  a  compressor  or  fan.   The source  legs  from the  various
                                   E-3
-------
Table  E-1.   PROCEDURE  TO  DESIGN  98 PERCENT  EFFICIENT  (60  fps,  300  Btu/scfJ
                      ELEVATED STEAM-ASSISTED SMOKELESS  FLARES
   Itm
                                                                    Value
I.  Haste gas flow rate, Qwg (scfm)a
Z,  Lower heating value of waste gas, LHVwg (Btu/scf)
3.  Teaperature of waste gas, Twg (°F)
4.  Holecular weight of waste gas, MWW<
5.  Ueight percent of hydrocarbons, wt. % HC
6.  Auxiliary natural gas flow rate,  Qng (scfm)b
     7.  Total  flare gas flow rate, Qfl_g (scfm)
     8.  Lower heating value of flare gas, LHVfl>g  (Btu/scf)

     9.  Teaperature of flare gas,  Tfl   (°F)C
    10.  Holecular weight of flare gas, MWfl   (Ib/lb-mole)
    II.   Calculated flare diameter, D calc. (in.)
    12.   Selected flare diameter, D(in.)e
    13.   Flare tip pressure drop.Ap (in. H20)
    1*.   Actual exit velocity, Vg  (fps)9
    IS.   Plane angle, 0h
    16.  Calculated flare height,  Hcalc> (ft)1

    17.  Selected flare height, H  (ft)
    18.  Safe pipeline length, L (ft)j
                                                                from Chapter 5
                                                                from Chapter 5
                                                                from Chapter 5
                                                                from Chapter 5
                                                                from Chapter 5
                                                                0, if  LHV   >300;
                                                            (300-LHV  )0
                                                                   "..     •  if LHVwn<30° Btu/scf
                                                                  {555}
                                                                                     wg
                                                            ^wg    ng
                                                            300,  if Qng > 0;
                                                            LHV if "ng - °
                                                            70,  if Qng = OorTwg-70;
                                                                «wg x "V + (17'4 X V
                                                                                        "ng'
                                                                                                      ,  if
                                                            (2.283 x 10"2)    IQfl _g /MWf1 _g (Tfl -------

































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-------
sources to the flare header were assumed to be 70 feet in length,8
while the length of pipelines to the flare was based on the horizontal
distance required to provide a tolerable and safe radiation level  for
continuous working (440 Btu/hr-ft^, including solar radiation).5
Piping and ducting were selected and costed as outlined in Section E.6.
     £.2.2  Flare Cost Estimation Procedure.  Flare purchase costs
were based on costs for diameters from 2 to 24 inches and heights  from
20 to 200 feet provided by National Air Oil Burner, Inc., (NAO) during
November 1982 and presented in Table E-2.5  A cost was also provided
for one additional case of 60 inch diameter and 40 feet height.1
These costs are October 1982 prices of self-supporting flares without
ladders and platforms for heights of 40 feet and less and of guyed
flares with ladders and platforms for heights of 50 feet and greater.
Flare purchase costs were estimated for the various regulatory alternatives
by either choosing the value provided for the required height and
diameter or using two correlations developed from the NAO data for
purchase cost as a function of height and diameter.  (One correlation
for heights of 40 feet and less, i.e., self-supporting flares and  one
for heights of 50 feet and greater, i.e., guyed flares.)  Purchase
costs of large diameter, 40-ft. high flares were approximated using a
curve developed from the NAO data  (see Figure E-l).  Purchase costs for
fluidic seals were approximated using a curve based on data provided by NAO?
(see Figure E-2).
     A retrofit installation factor of 2.65 (see Table 5-2) was used to
estimate installed flare costs.  Installed costs were put on a June
1980 basis using the following Chemical Engineering Plant Cost Indices:
the overall index for flares; the pipes, valves, and fittings index for
piping; and the fabricated equipment index for ducting.  Annualized
costs were calculated using the factors presented in Table 5-3.  The
flare cost estimation procedure is presented  in Table E-3.
E.3  THERMAL  INCINERATOR DESIGN AND COST ESTIMATION PROCEDURE
     Thermal  incinerator designs for costing  purposes were based on
heat and mass balances for combustion of the waste gas and any required
                               E-7
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     60,000—
     50,000-
     40,000—
     30..000-
     20,000-
     10,000-
                     10         20         30         40         50       60         70



                              Flare Tip Diameter  (in.)






                         «•

      Figure E-l.   Estimated Flare  Purchase Cost  for 40  ft Height
                                    E-9
-------
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                                                                            4J
                                                                             
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-------
auxiliary fuel, considering requirements of total  combustion air.
Costs of associated piping, ducting, fans, and stacks were also estimated.
E.3.1  Thermal Incinerator Design Procedure
     Designs of thermal  incineration systems for the various combinations
of waste gas streams were developed using a procedure based on heat
and mass balances and the characteristics of the waste gas in conjunction
with some engineering design assumptions.  In order to ensure a 98 percent
VOC destruction efficiency, thermal incinerators were designed to
maintain a 0.75 second residence time at 870°C (1600°F).9  The design
procedure is outlined in this section.
     Streams with low heat contents, which require auxiliary fuel  to
ensure combustion and sometimes require air dilution or fuel enrichment
to prevent an explosive hazard, are often able to utilize recovered
waste heat by preheating inlet air, fuel, and perhaps, waste gas.   The
                                              *• '•*•
design considerations for such streams are noted in the following
discussion, but the combustion calculations, etc. are not detailed
because all combined streams to thermal incinerators for polymers and
resins regulatory alternatives had sufficient waste gas heating values
to combust at 870°C  (1600°F) without preheating the input streams.
Therefore, only the design procedure for high heat content streams,
independently  able to sustain combustion at 870°C (1600°F), is detailed
in this section.
     The first  step  in the design  procedure was to calculate the
physical and  chemical characteristics affecting combustion of the
waste gas stream from the model plant characteristics given in Chapter 2,
using Table E-4.   In order to prevent an explosion hazard and satisfy
insurance requirements, dilution air was  added to any individual or
combined waste  stream with both a  lower heating value between 13 and
50 Btu/scf  at 0°C  (32°F)  (about 25 and  100  percent of the lower explosive
limit)  and  an oxygen concentration of 12  percent or  greater by volume.
Dilution air  was added to  reduce,the  lower  heating value of the stream
to below 13 Btu/scf.   (Adding dilution  air  is  a more conservative
assumption  than the  alternative of adding natural gas and is  probably
more realistic as  other  streams often have  enough heat  content to
sustain  the combustion of  the combined  stream  for the regulatory
alternative.)
                                E-ll
-------
      Table E-3.  CAPITAL AND ANNUAL OPERATING COST  ESTIMATION PROCEDURE FOR
                  STEAM-ASSISTED SMOKELESS FLARES
          Item
              Value
Capital Costs
Flare purchase cost, C'"fi
     (Oct. 1982 $)
Fluidic seal purchase cost,  C'"fi.s,
       (Oct.  1982  $)
Flare  system purchase cost,  C''fi
Select from Table E»2 if value given
or use equations:
  (3905.7) + (35.054) H x D + (900.36) D.
- (126.08)D2, for 20 < H < 40 ft and D <  8  in.
  (6275.6) + (224.10) H -f (12.782) H x D
+ (24.856)02, for 50 < H < 200 ft.
or from Figure E-2 if H = 40 ft and p >  8 in.
See footnote a.
C'"fl + C'"flc>s.
                                                                                '.< . ,,!! , „'. II i|l;|..|r '"If '"P"'	',,11
Flare  installed  cost, C'fj
      (Oct.  1982  $)
Total  installed  piping costs, C'p
      (Aug.  1978  $)
Total  installed  ducting costs, C'
      (Dec.  1977  $)
June 1980 Installed costs
        Piping0,  Cp
        Ductingd, Cd
        Flare6, Cf-| ^
        Total flare system cost,
          csys
C"fl x 2.65
Method  of  Appendix E.6
 Method  of  Appendix E.6
 C'p x 1.206
 C'd x 1.288
 C'
   fl  x 0.818
(Cp + Cd + Cfi)
                                         ;t JB;..	is	.•;-)
                                       E-12
                                                                                 ilf ', » |.J,:1! " J1. V .,"•!":
                                                                                ;	iii»^^^^^^       !
-------
             Table E-3.  CAPITAL AND ANNUAL OPERATING COST ESTIMATION PROCEDURE FOR
                         STEAM-ASSISTED SMOKELESS FLARES (Concluded)
                    Item
                Value
          Annualized Costs^

          Operating labor,

          Maintenance, Cm

          Utilities

                scfh
                scfh
          Cost natural gas, Cn n 1
                             1 1 • y •
          Cost steam,

          Taxes, admin. & insurance, Ctax

          Capital recovery, Ccr

          Total annualized,
  620 hr/yr x $18/hr = $11,160

  0.05 x Csys
                                                  80 scfh, for 2 < D < 8;
                                                 160 scfh, for 10 < D < 20;
                                                 240 scfh, for D = 24;
                                                 320 scfh, for D = 60.
 C(0.3272)(D in)2 -

            scfm
                                                                            cont] x 60
E
                         aux-
         scfh
53.45 C(Qn.g., pilot +

       (qscfh) purge)]
         n.g.

 3.296[QW _ (scfm) x MW x wt. % HC] cont,
         'y*                100%   fl.g.
 Cqv«; X 0.04
 0.1315 Cfi   + 0.1627 (Cp +

 CT  + Cm + Cn.g. -t- Cstm + Ccr + Ctax
                                             E-13
-------
Footnotes for Table E-3

apluidic seal is costed only if cost of purge gas without seal is greater
 than the annual i zed cost of the seal plus any purge gas required with
 the seal, i.e., taking the October 1982 purchase cost of a seal, Cf|.s.
 from Figure E-2 for D, if                                            '
53.45
60
      ™  {0.372
                                                              cont .
I
L
             (0.1315 + 0.05 + 0.04)
                                    $ capital
         x 0.818 Jun-  '80$ x 2.1  insta11ed   x  Cfl ". 1
                 Oct.  '80$        purchase       TI.S.J
            53.45 I/IL
                 scfh
              j A) .45
  or,  simplifying,

       if    1169|D(in.)|2  -  3154 /Q^f"1 \        >  (0.3805 x Cfl.s.)
                 L     J        \fl.g./cont.

       then  C'-'fT.s.  = Cfl.s.

             Note:   This condition  will be in error to the degree that1

            [0.45 D2 -  (QSCfm)       ] < 0
                         fl.g.  cont.
       Otherwise,
           C1"     = 0
              fl.s.
 bFor installation cost factor breakdown, see Table 5-2.

 ^Updated using Chemical Engineering Plant Cost pipes, valves and fittings
  index from August 1978 (273.1) to June 1980 (329.3).

 ^Updated using Chemical Engineering Plant Cost fabricated equipment index
  from December 1977 (226.2) to June 1980 (291.3),.

 eAdjusted using Chemical Engineering Plant Cost  Index from October 1982
  (317 estimated) to June 1980  (259.2).

 fpor annualized cost  factors,  see Table 5-3.

 QBased on vendor information for pilots without  energy  conservation  [
  (Reference 5).                                                      j

 "Ensures continuous flow of at least 1  fps for flare with  any  continuous
  flow not using fluidic seal:                         ,
                                ft'
                           x  (1 fps) x  (60  sec/min)
                      "•"   "  144  1

                  -  CQfl.g.  (scfm)]  cont. [ 60 min/hr

                                E-14
                                                           '.. ,	.',ii,, ...J	.i'lSW,,,! ;.  I,,'!!	:£«	E; "III;.; id
-------
Footnotes for Table E-3 (Concluded)

 Ensures sufficient continuous flow  per vendor information  for  flare
 with any continuous flow using a fluidic seal:

     0.45  scfh  x CD(in.)]2
             in
1<2CQn n (scfm).j x 60 min/hr x (t-j operating hours per year
      ."                             at stream combination i)
        CQpilot(scfh)
                                      x 8760 hr/yr
     x 520°R scf at 60°F  x 1,040 Btu (HHV) y     _ $5.98
       530°F scf at 70°R      scf at 60°F      (106 Btu (HHV)

               16
           „ (10° Btu)
           x 	g	
              10 BTU

JAssumes steam at 0.4 Ib/lb of hydrocarbon at maximum continuous flaring
 rate for 8600 hr/yr:

        Qcont  (scfra) X MWcont  x ( wt.% HC\       x 8600 hr/yr x 60 mi
                                  \   100% /cont.

        x (lb-mole/387 scf at 70°F)  x (0.4 Ib steam/lb HC)  x (1000 1b steam)
                                                             1000 (Ib steam)

        x $6.187(1000 Ib steam)

 or simplifying,
        3.2961Q  „ (scfm) x MW x  wt.% Hclcont.
             Lw-9-                 HHWjfl.g
                                  E-15
-------
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                                ,       ,      '•••  ' .("•  ii ',•  „ H"1  ,'•'•"  ,   :, , .    i  " 'i;,,,,:"!,
     The combustion products were then calculated using Table E-5
assuming 18 percent excess air for required'combustion air, but 0 percent
excess air for oxygen in the waste gas, i.e., oxygen thoroughly mixed
with VOC in waste gas.  The procedure wouldinclude a calculation of
auxiliary fuel requirements for streams (usually with heating values
less than 60 Btu/scf) unable to achieve stable combustion  at 870°C
(1600°F) or greater.  Natural gas was assumed as the auxiliary fuel as
it was noted by vendors as the primary fuel now being used by industry.
Natural gas requirements would be calculated using a heat  and mass
balance assuming a 10 percent heat loss in  the  incinerator.  Minimum
auxiliary fuel requirements for low  heating content streams would be
set at 5 Btu/scf to  ensure  flame  stability.10
                              I             ' •'».	' M.  	i."	;,;  '.  .  ;l	j ; '• ;  '   1  ' '.(.;•.
     The design procedure for streams  able  to  maintain  combustion at
870°C  (16b6°F) is  presented  in fable E-6.   Fuel was  added  for  flame
stability in  amounts that provided as  much  as  13  percent  of the  lower
heating value  of the waste  gas  for  streams  witn heating values  of
650 Btu/scf  or less.  For streams containing  morethan 650 Btu/scf,
                              |  i.. ,    •     - j	,Mi,i	i:"!,,1 fcii ' t -:."- • • •",!:!::'»	r :v"; • '  . :" i   '>
flame  stability fuel requirements were assumed to be  zero  since  coke
oven  gas,  which  sustains  a  stable flame,  containsonly about  590 Btu/scf.
In order to  prevent  damage  to incinerator construction materials,
quench air was added to reduce the combustion temperature to below the
incinerator  design temperature of 980 °C (1800 °F) for the cost curve
given by  IT Enviroscience.H
      The total flue gas was then calculated by summing the products of
combustion of the waste gas and natural  gas along with the dilution
 air.   The required combustion chamber volume was then calculated for a
 residence time of 0.75 sec, conservatively oversizing by  5 percent
 according to standard'industry practice.12  The design procedure
 assumed a minimum commercially available size of 1.01 m3  (35.7  ft3)
 based on vendor information13 and a maximum shop-assembled unit  size
 of 205 m3 (7,238 ft3).14                                             |
      The design procedure would  allow for  preiireating  of  combustion
 air, natural  gas, and when permitted by insurance guidelines, waste
 gas using a  recuperative heat exchanger  in order to  reduce the  natural
 gas required  to maintain a 870°C (1600°F)  combustion temperature.   If
                                    E-18
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                                                                           :syf
                                                                           'ft '
a plant had a use for it, heat could be recovered.  (In fact,  a waste
heat boiler can be used to generate steam, generally with a net cost
savings.)                                                            |
E.3.2  Thermal Incinerator Cost Estimation Procedure                 ;
     Thermal incinerator purchase costs for the calculated combustion
chamber volume were taken directly from Figure E-3, (Figure A-l in the
IT Enviroscience document, Reference 11).  A retrofit installation
cost factor of 5.29 (see Table 5-2) was used based on the Enviroscience
document.15  The installed cost of one 150-ft. duct to the incinerator
and its associated fan and stack were also taken directly from Figure
E-4 (Figure IV-15, curve 3 in the IT Enviroscience study16).  A minimum
                             I     .          . l:1 '  «,' ...I	'  : ..:>. ,•  '       }' I
cost of $70,000  (in December 1979) was assumed for waste gas streams
with flows below 500 scfm.  the costs of  pipingor ducting from the
process sources to the 150-ft. duct costed above were estimated as for
flares.  Installed costs were put on a June 1980 basis using the
following Chemical Engineering Plant Cost Indices:  the overall index
for thermal incinerators; the pipes, valves, and fittings index for
piping; and the  fabricated equipment index for ducts,  fans, and stacks.
Annualized costs were  calculated  using the factors  in  Table 5-3.  The
                             I                !   , !•, , • I '              '  1 ' ' '
electricity required was calculated assuming a 6-inch  ^0 pressure
drop across the  system and a blower efficiency of 60 percent.  The
cost calculation procedure is given in Table E-7.                    ;
                             :    ,    .        ,   Vii   ,. I  ••.;     '•   "i, . i   .1
E.4  CATALYTIC INCINERATOR DESIGN AND COST ESTIMATION  PROCEDURE      :
     Catalytic incinerators  are generally cost effective VOC control
devices  for low  concentration streams.   The catalyst increases the
chemical  rate of oxidation allowing the  reaction  to proceed at  a  lower
energy level  (temperature) and thus requiring a smaller  oxidation
chamber,  less expensive  materials,  and much less  auxiliary  fuel
 (especially  for  low  concentration streams) than required  by a  thermal
incinerator.   The  primary determinant  of catalytic  incinerator capital
cost is  volumetric flow  rate.  Annual  operating costs  are  dependent  on
emission rates,  molecular weights,  VOC concentration,  and  temperature.
Catalytic  incineration in conjunction  with  a  recuperative  heat  exchanger
     **                      i             ..'.'•,'-,;  ';„ .;:     : -    ;   ;  |
can reduce overall fuel  requirements.
.jj:	r.
                                   E-22
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                                                 E-23
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                                                                         E-24
-------
        TABLE E-7.   CAPITAL  AND  ANNUAL  OPERATING  COST  ESTIMATES  FOR
            RETROFIT THERMAL INCINERATORS  WITHOUT HEAT RECOVERY
        ITEM
               VALUE
Capital Costs

Combustion Chamber

 Purchase cost
 Installed cost
 Installed cost, June 1980a

Piping & Ducting (from sources
 to main incinerator duct)

 Installed cost

 Installed cost, June 1980b
Ducts, Fans & Stacks (from
 main duct to incinerator
 and from incinerator to
 atmosphere)

 Installed cost0
  Installed cost, June 1980d

Total  Installed Cost, June 1980



Annualized Costs6

  Operating labor
  Maintenance material & labor
  Utilities

    natural gas


    electricity^

  Capital  recovery"

  Taxes,  administration &  insurance

  Total Annualized Cost
from Figure E-3 for Vcc
purchase cost x 5.29
installed cost x 1.047
see Section E.6 for Qw.g. (scfm)

installed cost x 1.206 for piping
installed cost x 1.288 for ducting
from Figure E-4 for Qw q ;
  use $70,000 minimum

installed cost x 1.064

sum of combustion chamber,
piping & ducting, and ducts,
fans, & stacks
 1200 hr/yr x $18/hr = $21,600
 0.05 x total installed cost


 (5.245 x 10'4)  (% aux) x LHVW _
 x Qw.g. (lb/hr)            W*9'

 (0.4610) x Qf>g> (scfm)

 0.1627 x total  installed cost

 0.04 x total installed cost

 operating labor + maintenance
  + utilities + capital recovery
  + taxes, administration &
    insurance
                                E-25
-------
Footnotes for Table E-7

aUpdated using Chemical Engineering Plant Cost Index from December 1979
 (247.6) to June 1980  (259.2).

bPiping updated using  Chemical Engineering Plant Cost pipes, valves,
 and fittings index from August 1978 (273.1) to June 1980 (329.3).
 Ducting updated using Chemical Engineering Plant Cost fabricated
 equipment index from  December 1977 (226.2) to June 1980 (291.3).

cFrom Figure E-4 for no heat recovery from Enviroscience (Reference  16),
 which assumed 150-ft  of round steel inlet ductworkwith four ells,
 one expansion joint,  and one damper with actuator;, and costed according
 to the CARD Manual (Reference 17).  Fans were assumed for  both waste
 gas and combustion air using the ratios developed for a "typical
 hydrocarbon" and various estimated pressure drops and were costed
 using the Richardson  Rapid System  (Reference 18).  Stack costs were
 estimated by Enviroscience based on cost data received from one
 thermal oxidizer vendor.

 Although these Enviroscience estimates were developed for  lower
 heating value waste gases using a  "typical hydrocarbon" and no dilution
 to limit combustion temperature, the costs were used directly because
 Enviroscience found variations in  duct, etc., design to causeonly
 small  variations in total system cost.  Also, since the duct,fan,
 and stack costs are based on different flow rates  (waste gas, combustion
 air and waste gas, and flue  gas, respectively) 'thecosts can not  be
 separated to be adjusted individually.

^Updated using Chemical Engineering Plant Cost fabricated equipment
 index  from  December 1979  (273.7) to June 1980 (291.3).

eCost factors presented in Table  5-3.
               .  '         ''. f •.  I      .    '  •':• ,'N Vil;M>!": , | '.''• ,':'• •;••'  - ;1Ls:; ', \'\"
f[(% aux) x  LHVw>g/20,660 Btu/lbn>g;]  (100  lbn>g/100  lbw>gj x Q    '(Ib/hr)  x

 (100 lbw.gj/100(lbw.gj  x  (8000  hr/yr)  x  (lb-mole/17.4  lbn.g.) x

 (379 scf  at 60°F/lb-mole) x  (1040  Btu(HHV)/scf  at WF)  x  $5.98/10^
                  ; _  ,.  •   .     I     ,       ,'.'( ; '  ;; - :(•,!.•(' ]  • [ ,:;  •.'•• t,,;  ;,,,„;;.:',; ';•. !
 Btu  (HHV) x (106  Btu)/106  (Btu).

SElectricity =  (6  in.  H20 pressure drop)  x  Qf.g.  (scfm)  x  (8000  hrs/yr)

 X (0.7457  kW/hp)  x (5.204 Ib/ft2/in.  H20)  *   [(60 sec/min) x (550 ft-lb/

 sec/hp)  x  (0.6  kW blower/1  kW electric)  x  $0.049/kwh].

n!0 percent  interest   (before  taxes) and 10  yr.  life.
Hi!"
                                E-26
-------
E.4.1  Catalytic Incinerator Design Procedure
     The basic equipment components of a catalytic incinerator include
a blower, burner, mixing chamber, catalyst bed, an optional  heat
exchanger, stack, controls, instrumentation, and control  panels.  The
burner is used to preheat the gas to catalyst temperature.  There is
essentially no fume retention requirement.  The preheat temperature is
determined by the VOC content of gas, the VOC destruction efficiency,
and the type and amount of catalyst required.  A sufficient  amount of
air must be available in the gas or be supplied to the preheater for
VOC combustion.  (All the gas streams for which catalytic incinerator
control system costs were developed are dilute enough in air and
therefore require no additional combustion air.)  The VOC components
contained in the gas streams include ethylene, n-hexane,  and other
easily oxidizable components.  These VOC components have catalytic
ignition temperatures below 315°C (600°F).  The catalyst bed outlet
temperature is determined by gas VOC content.  Catalysts can be operated
up to a temperature of 700°C (1,300°F).  However, continuous use of
the catalyst at this high temperature may cause accelerated thermal
aging due to recrystallization.
     The catalyst bed size required depends upon the type of catalyst
used and the VOC destruction efficiency desired.  About 1.5 ft^ of
catalyst for 1,000 scfm is required for 90 percent control efficiency
and 2.25 ft^ is  required for 98 percent control efficiency.^  As
discussed earlier many factors influence the catalyst life.   Typically
the catalyst may loose its effectiveness gradually over a period of
2 to 10 years.   In this report the catalyst is assumed to be replaced
every 3 years.
     Heat exchanger  requirements are determined by gas inlet temperature
and preheater temperature.  A minimum practical heat exchanger efficiency
is about 30 percent.  Gas temperature, preheater temperature, gas dew
point temperature and gas VOC content determine the maximum feasible
heat exchanger  efficiency.  A maximum heat exchanger efficiency of
65 percent was  assumed for this analysis.  The procedure used to calculate
fuel requirements is presented in Table E-8.   Estimated fuel requirements
and costs are based  on using natural gas, although either oil (No. 1
or 2) or gas can be  used.  Fuel  requirements are drastically reduced
                                  E-27
-------
                                                      K '-ll';;' K.	-I';*'''',,,:*	'
when a heat exchanger  is  used.   Total  heat requirements are based on a
preheat temperature  of 600°F.   A stack is used to vent flue gas to the
                               I  j,  ' ' ,   '  " '! :„",  ':•'• -/ilk! '.'U'.;! !	'> .if I':;.-  -n ., •"  I'-1
atmosphere.
E.4.2  Catalytic  Incinerator Cost Estimation Procedure
     The capital  cost  of  a catalytic incinerator system is usually
based  on gas  volume  flow  rate at standard conditions.  For  catalytic
incineration, 70°F and 1  afm (0 p'sig)  were taken as standard  conditions.
The  operating costs  are determined from the gas  flow  rate and other
conditions  such as gas VOC content and temperature.Table  E-9 presents
       .  i                      ,!•"•:,•,  '	  	i-;.' •	"»V'"	l!"!"1 ••! i' i ±~>f' ' ' *•'•'   •' I ' '•	-  '
the  basic  gas parameters  required for  estimating system  costs.
     As  noted earlier, equipment components of a catalytic  incineration
system include blower, preheater with  a  burner,  mixing  chamber, catalyst
bed, an  optional  heat exchanger, stack,  controls,  and internal ducting
including  bypass.  Calculations  for capital cost estimates  are based
on equipment purchase costs obtained  from vendors 19,20,21 and application
of direct  and indirect cost factors.   Table E-10 presents third quarter
1982 purchase costs of catalyst incinerator systems with and without
heat exchangers  for sizes  from  1,000  scfm to  50,000 scfm.  The cost
 data are based on carbon  steel  for incinerator systems and stainless
 steel  for heat exchangers.  The heat  exchanger cosis are based on
 65 percent heat  recovery.   Catalytic  incinerator systems of  gas volumes
 higher than  50,000  scfm  can be estimated by considering two  equal
 volume units in  the system.  A minimum availableunit size of  500  scfm
 was assumed.22'23  The installed cost of this minimum size unit  (which can
 be  used without  addition of gas or air for stream flows greater  than
 about 150  scfm23) was estimated to be $53,000 (June  1980).   The  heat'
 exchangers  for small  size systems would be costly and may  not be practical
 Table 5-2  presents  the direct and indirect installation cost component
 factors used for estimating capital costs of catalytic  incinerator
 systems.   The geometric  mean of the two  vendor  estimates for each  flow
 rate  was  multiplied by the ratio of total  installed  costs  to equipment
 purchase  costs of 1.82 developed for  a  skid-mounted  catalytic incinerator.
 Actual  direct and indirect cost  factors  depend  upon  the plant specific
 conditions and may vary with system  sizes.
       Since the  equipment  purchase  cost  presented  in Table  E-10
  represents the third quarter of  1982, the cost  data was adjusted to
                                     E-28
                                •ii , .v,i	i	y.	iJr	./(!,	Jj''fei&&BSi:H!4:.VA iik.\\*.	^iii•:«.';i(j.
                                                                            I:'1 i:11!"!1"'' "!!:l<:*"1"I:I!
                                                                            .••'•r;^.^M
                                                                            , 1!l|i"li'ii: S:ll! i'S
                                                                            " ii	Bi' rl^'.il
•	'i!»5, !,:«!*'
 'i; ""»! '.S! iK-'t
-------
     Table E-8.  OPERATING PARAMETERS AND FUEL REQUIREMENTS
                     OF CATALYTIC INCINERATOR SYSTEMS
     Item
  Source of information or calculation
Waste Gas Parameters

(1)  Flow rate (0.2), scfm

(2)  Amount of air present in
     the gas, scfm
(3)  Amount of air required
     for combustion at 20%
     excess, scfm

(4)  Net amount of additional
     air required (0.3), scfm

(5)  Total  amount of gas to be
     treated (04), scfm

(6)  Waste gas Temperature at
     the inlet of PHRb, °F

(7)  Waste gas temperature at
     preheater outlet or
     catalyst bed inlet, °F

(8)  Temperature rise in the
     catalyst bed, °F

(9)  Flue gas temperature at
     catalyst bed outlet, °F

(10) Minimum possible temperature
     of flue gas at PHR outlet, °F

(11) PHR efficiency at maximum
     possible heat recovery**, %
(12) PHR design efficiency, %
      From Table E-9

  0, if the waste gas contains VOC and
  nitrogen or other inert gas; and
  [(1 - volume percent VOC) * (volume
  percent VOC)] x VOC volume flow (O^)
  scfm, if the waste gas contains VOC
  and air

  See footnote a.
 Item (3) - Item (2); and 0 if
 [Item (3) - Item (2)] is negative

 Item (1) + Item (4)
 From Table E-9
 600°F
(25°F/1% LEL) x (%LEL from Table E-9)
Item (7) + Item (8)
See footnote C.
[Item (1) x (Item (7) - 25°F -
Item (6))] * [Item (5) x (Item (9) -
Item (6))]e

See footnote f
                                    E-29
-------
                                    I          ,   '.   '   "I  , I  ' • 11,

          Table  E-8.   OPERATING  PARAMETERS AND FUEL  REQUIREMENTS
                       OF  CATALYTIC  INCINERATOR SYSTEM  (concluded^
     Item
                                       Source  of  information or calculation
(13)  Waste gas temperature at
     PHR outlet,°F

(14)  Amount of heat required by
     preheater at additional 10%
     for auxiliary, Btu/min

(15)  Amount of heat required
     for preheater and auxiliary
     fuel, 106 Btu/h

(16)  Amount of natural gas
     required per year, 106 cfm
0.65 [Item (9) - Item (6)3 + Item (6)
Item (5) x [Item (7) - Item (13)] x
[Gas specific heatS, Btu/scf, °F] x
[Item (14) x 60 minutes/hour] x (lO%)n
x (106 Btu)/106 Btu
[Item  (14) x (8,tiOO x 60) minutes/year]
x 10-3 *  (1,040 Btu/cfm)
 „„ volume basis (scfm/scfm):  11.45 for methane, 20.02 for ethane, 28.58 for
 propane, 54.31 for hexane, 17.15 for ethylene, and 45.73 for pentane.
 Values taken from p. 6-2 in Steam (Reference 24) for 100% total air and
 multiplied by 1.2 for 120% total air or 20% excess air.

bPrimary heat recovery unit.

CHeat  exchanger should be designed for  at least 50°F above the  gas dew point.

dThe heat exchanger will be designed for 25°F lower thanthe preheater
 temperature  so as to  not cause  changes in  catalyst bed outlet  temperature.

^Though the heat  recovery to the temperature level of inlet gas is the
 maximum heat efficiency possible, in  some  cases  this may not be possible
 due to gas dew point  condition.

fCost  estimates  are  based  on  calculated maximum  possible heat recovery
 up to an  upper  limit  of 65  percent  heat  recovery.

9Gas specific heat varies  with composition  and. temperature.  Useci  6.019  Btu/ft3°F
 based on  average specific  heat of air for  calculation  purpose.

 ^Auxiliary fuel  requirement  is assumed to be  10  percent  of  total.
                                          E-30
                                                                                     	i
-------
        TABLE E-9.  GAS PARAMETERS USED FOR ESTIMATING CAPITAL AND
               OPERATING COSTS OF CATALYTIC INCINERATORS3
          ITEM
                  VALUE
Stream identification

Stream conditions
  Temperature,°F
  Pressure, psig
  VOC content:
     Emission factor, kg/Mg
       of product
     Weight % of total gas
     Mass flow rate, kg/h
           Ib/h

    .Organic constituents, wt %
     Average mol. wt. (M]_), IDS
     Volume flow (Qi), scfm
     Heat content
       Btu/scf
  Total gas:
     Constituents
     Mass flow rate, Ib/h
     Molecular weight (M2)
     Volume flow (0.2)5 scfm
     Air volume flow rate, scfm
     VOC concentration (A), %
       of LEL

     Heat content
       Btu/total scf
Identify the vent and the polymer
industry from Chapters 2 and 5
(Emission factor, E, kg/Mg) x 1000 Mg/Gg
(Plant production rate, P, Gg/yr) *
(8,000 h/yr)
(kg/h) x (2.205 Ib/kg)
(VOC mass rate, Ib/h) * (60 min./h) 4
(Molecular weight (M^), Ibs/lb mole) x
 385 scf/lb-mole at 68°F) = 1.768
(174.273)(2.521NC + NH)
VOC, air and others
(VOC rate, Ib/h) * (wt% of VOC in
gas, Wj/100%)
Gas mass rate, Ib/h) •* (60 min/h) *
(Gas molecular weight (M2), Ib/lb mole)
x (385 ft6 /Ib mole) = 1.768
(Total gas flow (Q2), scfm) - (VOC volume
flow (Qi), scfm)

(100) [(Volume flow of VOC, scfm) *
(Volume flow of air, scfm] * LEL"
From Chapter 5e
                                 E-31
-------
Footnotes for Table E-9

^Obtain gas parameters from Chapter.2  of  the CTG,  and Chapter 3 of the
 background information document  for the  polymer manufacturing NSPS,
 except those to be calculated.
          •                     '           ..'•.• iiiii, i. i«  '
bCalculate using weight percent  values of VOC components.
                                       1   '•"    •••	""•'•'' ' ! •
cif the VOC heating value is  not available, calculate it using heat of
 combustion values of 14,093  Btu/lb from carbon converted to C02 and
 51,623 Btu/lb  from hydrogen  converted to water.  Nc and NH denote  number
 carbon and hydrogen  atoms in VOC.

dlower explosion  levels  of ethylene, hexane, methanol, propane, butane,
 and  pentane  are  3.1, 1.32, 7.3, and  2.5, 1.9,.and  1.4, respectively.

eiotal gas  heat content averages 50 Btu/scf  at  100 percent LEL.
                                     E-32
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                                                                        '"•  " !R It/if:	!![;'•¥'	IW	ii
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                                                                        '
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Figure E-5.   Installed Capital Costs for Catalytic  Incinerators
             With and Without Heat Recovery
                        E-35
-------
          Table £-11.   CAPITAL  AND  OPERATIONG  COST'ESTIMATION" FOR
                             CATALYTIC  INCINERATOR  SYSTEMS
                                                                                  il; WJ
                                                                                    '
           Item
                                                      Value
Capital  Costs
    Incineration system
      Installed cost, June 1980

      Installed retrofit cost, June 1980

    Piping & ducting (from sources
      to main incinerator duct)

      Installed cost
      Installed cost, June 1980a
    Ducts, fans & stacks  (from main duct
      to  incinerator and  from incinerator
      to  atmosphere)

      Installed costb
       Installed  cost, June  1980C

     Total  Installed  Cost, June 1980


 Annualized Costs
   Direct costs
     Operating labor
     Maintenance material  and
       labor

     Catalyst requirement
     Utilities:
       Fuel (natural gas)
From Figure E~5

Installed cost x 1.18,  from Table 5-2
See Section E-6 for source flow
rates, scfm.
                              i

Installed cost x 1.206 for piping
Installed cost x 1.288 for ducting
                  '
 From Figure E-4 for waste gas flow
 (Q2),  scfm; use $70,000 minimum

 Installed cost x 1.064

 Sum of incineration systems,
 piping & ducting,  and ducts,
 fans,  & stacks


 $11,200 for  systems with  no  heat
 recovery; and  $16,700 for  systems
 with  heat  recovery
 (0.05) x (Total  installed capital
 cost, $ from Figure E-5)
                              I
 $2.7 x (Total  gas volume  flow|(Q4)a
 scfm, item 5 from Table E-8)  =
 ($2.7 x Q4)

 ($6.22/103ft3) x (Amount of natural
 gasrequired, 103ft3, Item 16 of
 Table E-8)e
                                     E-36
-------
          Table E-ll.  CAPITAL AND OPERATIONG COST ESTIMATION FOR
                       CATALYTIC INCINERATOR SYSTEMS (Concluded)
           Item
         Value
      Electricity
Indirect Costs
    Capital recovery
    Taxes, insurance and
      administrative charges
Total Annual!zed Costs
($0.312/scfm) x (Total gas volume
flow rate (0,4), scfm, Item 5 from
Table E-8) for units with no heat
recovery; and ($0.78/scfm) x (Total
gas volume flow rate (04), scfm,
Item 5 from Table E-8) for units with
heat recovery

(0.1627) x (Total installed capital
cost, $ from Figure E-5)

(0.04) x (Total installed capital
cost, administrative charges
$ from Figure E-5)

Sum of total  direct costs and
total indirect costs
aUpdated using Chemical Engineering Plant Cost Index from December 1979
 (247.6) to June 1980 (259.2).

^Piping updated using Chemical Engineering Plant Cost pipes, valves, and
 fittings index from August 1978 (273.1) to June 1980 (329.3).  Ducting
 updated using Chemical Engineering Plant Cost fabricated equipment index
 from December 1977 (226o2) to June 1980 (291.3).

cSee footnote c, Table E-7 for discussion on application of these costs
 developed by Enviroscience (Reference 25).

^Updated using Chemical Engineering Plant Cost fabricated equipment index
 from December 1979 (273.7) to June 1980 (291.3).

eTotal gas flow including waste gas and additional  combustion air.
                                     E-37
-------
$0.335/scfm for units with no heat recovery (i.e., for 4 in.  1^0 pressure
drop) and $0.838/scfm for units with heat recovery (i.e., for 10 in.  ^0
                                                                     i
pressure drop).
E.5  SURFACE CONDENSER DESIGN AND COST ESTIMATION PROCEDURE
     This section presents the details of the procedure used for
sizing and estimating the costs of condenser systems applied to the
gaseous streams from the continuous process polystyrene model plant.
Two types of condensers are in use in the industry:  surface condensers
in which the coolant does not contact the gas or condensate; and
contact condensers in which coolant, gas, and condensate are intimately
mixed,               "         '    '        , '("'	'' '",','    	   j
     Surface condensers were evaluated for the following two streams
from the polystyrene model plant:  the styrene condenser vent and the
styrene recovery unit condenser vent.  These streams consist of styrene
and steam, which are immiscible, or of styrene inair, a non-condensable.
The nature of  components  present in the  gas  stream determines the
method of condensation:   isothermal or non-isothermal.  The condensation
method for streams containing  either  a pure  component  or a mixture of
two immiscible components is isothermal.   In the  isothermal condensation
                              i u '           •,      "i	','i.r  ,i ,  .• ,, .  i«  • >• -,	  , 1 ,  1 •	
of two immiscible  components,  such  as styrene and  steam, the components
condense at the saturation temperature and yieldtwo  Immiscible liquid
condensates.   The  saturation temperature is  reached when the vapor
pressure of the components equals the total  pressure of  the  system.
The  entire  amount  of vapors  can be  condensed by  isothermal  condensation.
Once  the condensation temperature  is  determined,  the total  heat load  is
calculated  and the corresponding heat exchanger  system size  is  estimated.
The  condensation of  styrene  mixed  with a non-condensable,  such  as  air,  _
can  be considered isothermal  if the temperatureof one fluid is nearly
constant.   The analysis shows  that  the condenser coolant tempearture  is
 nearly constant for the combined material recovery vent stream from  the
continuous  polystyrene model  plant.  The condensation of styrene  in
 air, nevertheless, is accomplished less  readily, and thus more expensively,
 than the condensation of styrene in steam.
      The following procedures and assumptions were used in evaluating
                                                 	•;  ; ; "    ,,';,'  ','     i     ;„
 the isothermal condensation  systems for  the two streams containing
                                   E-38
-------
(1) styrene in steam and (2)  sytrene in air from the continuous polystyrene
model plant.
E.5.1  Surface Condenser Design
     The condenser system evaluated consists of a shell  and tube heat
exchanger with the hot fluid in the shell  side and the cold fluid in
the tube side.  The system condensation temperature is determined from
the total pressure of the gas and vapor pressure data for styrene and
steam and sytrene in air.  As the vapor pressure data are not readily
available, the condensation temperature is estimated for styrene in
steam by trial-and-error, and for styrene in air by a regression equation
of available data points'^ using the Clausius Clapeyron equation which
relates the stream pressures to the temperatures.  The total pressure
of the stream is equal to the vapor pressures of individual components
at the condensation temperature.  Once the condensation temperature is
known, the total heat load of the condenser is determined from the
latent heat contents of styrene and steam and, for styrene in air, from
the latent heat content of the condensed sytrene and the sensible heat
changes of styrene and air.  Table E-12 shows the procedure for calculating
the heat load of a condensation system for styrene in air.  The design
requirements of the condensation system are then determined based on
the heat load and stream characteristics.   The coolant is selected
based on the condensation temperature.  The condenser system is sized
based on the total heat load and the overall heat transfer coefficient
which is established from individual heat transfer coefficients of the
gas stream and the coolant.  An accurate estimate of individual coefficients
can be made using such data as viscosity and thermal conductivity of
the gas and coolant and the standard sizes of shell and tube systems to
be used.
     For styrene in steam, no detailed calculations were made to determine
the individual and overall heat transfer coefficients.  Since the
streams under consideration contain low amounts of styrene, the overall
heat transfer coefficient is estimated based on published data for
steam.
     For styrene-in-air, refrigerated condenser systems were designed
according to procedures for calculating shell side^S and tube side2^
heat transfer coefficients and according to condenser^ and refrigerant31>32
                                  E-39
-------
                Table E-12.  PROCEDURE TO CALCULATE HEAT LOAD
                 OF A CONDENSATION SYSTEM FOR STYRENE I
         Item
Heat exchanger type
Source identification


Source production capacity
  (CAP), Gg/yr

Source emission  factor  (E),
  kg  VOC/Mg  product

Desired emission reduction,
  (%  Red'n), %
 Gas stream condition


 Partial pressure of styrene
  at inlet (P1n)

 Composition of gas stream
  at inlet;

 Styrene mass flowrate
  (Ms), lb/hrd;

 Gas stream volumetric
  flowrate  (V),  acfrn6

 Gas stream mass  flowrate
  (W),  lb/hrf

 Partial  pressure of styrene
  at outlet (P0ut)>  mm H9

 Temperature required for  reduction
  Temperature required for reduction
   (Tout). "F
  Latent heat change of styrene
   Btu/hrh
                                                            Value
Shell and tube heat exchanger
 with hot fluid in the shell
 side and the cold fluid in the
 tube side
Identify the polymer industry and
 the vent from Chapters 2 and 5
     • H.1. r . i     ;;	'  ; i ..];•. •   • :	-

From model plant in Chapter 2
 From model  plant  in Chapter 2
 96.1%  at  3.09  kg  VOC/Mg  of  product
 40% at 0.2  kg  VOC/Mg  of  product

 Assume^ saturated styrene in
  air at 80°F,  latm.
 7.952 mm Hg
  .      ft3 styrene/ft3 gasb;
 0.002764 Ib styrene/ft* gasc
 0.2756 x CAP x E
 361.79 x
  4.415  x V
  100-%Red'n    x  7.952 mm Hg
     100
   4847.95 * [18,2440 - In (Pout)]
  (1.8 x T'out) - 459.67


  166.36 x W x (% Red'n)
                                       E-40
                                                                                  111! I 111
-------
                Table E-12.  PROCEDURE TO CALCULATE HEAT LOAD
                OF A CONDENSATION SYSTEM FOR STYRENE IN AIR (Concluded)
         Item
                                                        Value
Average (bulk) gas temperature (Tb),°F

Density of air (pair), Ib/ft

Specific heat of air((cD)air),
 Btu/lb-°F             '

Sensible heat change of air (Qa-jr)»
 Btu/hr
 Specific heat of styrene ((CD)   ),
  Btu/lb-°F
(80 + Tout)  * 2

1 * [(0.002517 x Tb) + 1.157]


From API Report 44k

V x pair x (c) .  x (80-Tout) x

  60 min/hr
                                                          p  .        -out
                                                            ai r
                                             From API Report 441
Sensible heat change of styrene (Q'sty)
 Btu/hr
Total design heat load (Qt0t)> Btu/hrm
                                                   (cp)    x (80-Tout)
                                                       sty
                                             1.2 (Qsty+Qair+ Q'sty)
Calculated from Clausius Clapeyron curve fit

        (In p = m     \   + b)  of styrene vapor pressure versus


 temperature data given on p. 3-59 of the Chemical Engineers' Handbook
 (Reference 26) for 80°F (see temperature required for reduction).

bVolume fraction of styrene = 7.952 mm Hg = 0.01046 ft3 styrene/ft3 gas.
                                760 mm Hg

cAssuming ideal gas:

    V  _ RT =  1545 ft lbf/lb.m - °R x0540^°R  = oqd Tjft3/u
    n"  " F"     14.7 lbf/in.2 x 144inz/ftz      «•»•"   /'L

    styrene content (Ib/ft3 gas) -

        0.01046 ft3 styrene x  Ib-mole    x 104.14 Ib styrene
             "•° gas          394.13 ft3        1 b-mole
                                   E-41
-------
Footnotes for Table E-12 (Concluded)


dCAP Gg product/yr x 1000 Mg/Gg x E kg VOC/Mg product
 	8000 hr/yr x 0.4b3b Kg/ID

eCAP Gg product/yr x 1000 Mg/Gq x E kg VOC/Mg procjuct  = 92.26 x CAP x E
 	'     8000 hr/yr x 0.4b^6 kg/lb x u.uuz/64

fV, acfm @ 80°F x  ^29 Ib/lb-mole x 60 min/nr
 —394.13 acT/lb-mole e birt-

gsolving Clausius  Clapeyron curve fit of  styrene  vapor pressure data
  (?2 1  0.99995) referred to in  footnote a for temperature.

hsiope, m,  of Clausius  Clapeyron  curve fit = - X/R

  latent heat of  styrene,X= -m x  R =              ptu-lb-mole

       4847  95 (°K) x 1.9853 cal/g-mole-°K x 1.8  cal/g-mole
       	'-—	104.14 Ib/lb-mole

 1T 6R = T oc + 273.15 = 5. (T,dF-32)  + 273.15 = 0..5556 T,°F + 255.37
                         9
    an ideal  gas (PV = mRT/MW) ,  £1 =
                                P2
at 0°C (ChE Hndbk, p. 3-72)


   Pair @ T,°K = 0.0808 X
   'all    "
                                                   ;' Pair = 0.08081
                               T,6K      0.5556 x f,°F + 255.37

  k(cp)  .   =   0.796 (cp)N2  +  0.231  (cp)02,  where  (cp)N2 and  (cp)02 are

                                                     !  i       , "       i

   specific heats of nitrogen and air,  respectively,  available  by
   interpolation from API Report 44,  p. 652 (Reference  27).

  V(c )    vs T,°F, values are available for interpolation on p.  682
     P sty

   of API Report 44 (Reference 27).

  ^including 20% safety margin.
                                     E-42
                                                                            	i	      I
-------
characteristics given primarily in the Chemical  Engineers'  Handbook  and
consistent with the 8-ft. long condenser with 1-inch outside diameter  tubes
assumed by Enviroscience33 for cost estimation purposes.   Then the total
heat transfer area is calculated from the known values of total  heat
loads and overall heat transfer coefficient using Fourier's general
equation.  A tabular procedure for calculating heat exchanger size is
presented in Table E-13 for styrene in steam and in Table E-14 for
styrene in air.
E.5.2  Surface Condenser Cost Estimation Procedure
For styrene in steam, the heat exchanger costs for each stream were
obtained from vendors.36>37,58 por styrene in air, condensation system costs
were based on IT Environscience3^ as well as vendor information.
A retrofit installation factor of 1.48 (See Table 5-2) was used to
estimate installed condenser costs for condensers of 20 ft2 or less  and
2.58 for condensers 125 ft2 or greater.  No additional piping was costed
for condensers with less than 20 ft2 of heat transfer area because
the condenser unit is so small ,( 1-2 ft. diameter) that it should
be able to be installed adjacent to the source.  For condensers with
heat transfer areas of 125 ft2 or greater, piping was costed using the
procedures described in Section E-6.  Table E-15 presents the estimated
total capital and annual operating costs for the condenser system of 20
ft2 heat transfer area for styrene in steam.  Table E-16 presents the
procedure for estimating capital and annual operating costs for condensation
systems for  styrene  in air.
E.6  PIPING  AND  DUCTING  DESIGN AND COST ESTIMATION  PROCEDURE
     Control costs for flare and incinerator systems included costs of
piping or ducting to convey the waste gases  (vent streams) from the
source to a  pipeline via a source leg and through a pipeline to the
control device.  All vent  streams were assumed to have sufficient
pressure to  reach the control device.   (A fan is included on the duct,
fan, and stack  system of the  incinerators.)
E.6.1  Piping  and Ducting  Design Procedure
     The  pipe  or duct diameter for each waste gas stream (individual
or combined) was determined by the procedure given  in Table E-17.  For
flows  less  than  700  scfm,  an  economic pipe diameter was calculated
                                   E-43
-------
Table  E-13    PROCEDURE TO  CALCULATE  HEAT TRANSFER  AREA OF  AN
                          ISOTHERMAL CONDENSER SYSTEM
              Item
         Heat exchanger type
         Gas stream condition
           (including temperature
           (T,)°F pressure  (PJ..
           psig, and composition)
         Condensation temperature
           (T2),°F
         Total  heat load (H),  Btu/h
          Coolant used0
          Temperature, rise of
            coolant,  (AT),°F
          Coolant outlet temperature
            (T3),8F
          Log mean temperature
            difference (LMTD),°F
          Heat  transfer coefficient (U)
           Heat transfer area (A), ft
       Value
  Cocurrent shel1  andtube heat
  exchanger with the hot fluid
  in the shell side and the cold
  fluid in the tube side
  Obtain from Chapters 2 and  5
   water at 85°F,  25 gpm
   H Btu/h * [(25  gpm  x 500 Ib/h/gpm) x
   (1 Btu/lb°F)]
   85°F + AT
    [(TrT3) - (T2 - 85)3  *  in  C(TrT3)/(T2-85)]
    240 Btu/h  ft°F
    (H)/U(LMTD)
                              in P/Pr
(X/R)  (1/T0 - 1/T)
            and R is universal gas constant = 1.99 cal/g mole K.
            The same equation can be rearranged to eliminate X and R:
                              ln(P/P)
             (970 3 Btu/lb  steam) and Ib/hr steam  in stream.
            cF1xed amount of 25 gpm is used in order to maintain turbulent flow.


             s?eL andl6% |tyrene) of P"« «««l;«tfSR2sgJ
             1,000 Btu/h ftz°F for steam and 35 Btu/h  ft  F tor
             following relationship:
                                          E-44
-------
                  Table E-14.  PROCEDURES TO CALCULATE HEAT TRANSFER
                    AREA OF A CONDENSATION SYSTEM OF STYRENE IN AIR
Heat exchanger configuration
Source Identification
Coolant temperature, TC,°F
Shell-side heat transfer
 Coefficient (h0),      Btu
                    hr-ft2 - °F
Try 8" shell with 17 1-inch o.d.,
16 gage, 8-feet long brass tubes
on 1-1/4" square pitch9

Identify the polymer industry and
the vent from Chapters 2 and 5

Tout-10, rounded to next lower

multiple of 5.

Calculate using procedure in
Chemical Engineers' Handbook,
pp. 10-25 thru 10-28
(Reference 28)b
Coolant
Tube-side Reynold's Number (

Tube-side heat transfer
 Coefficient (h0),      Btu
Select chilled water at Tc > 35°F;

for Tout > 45°F; ethylene glycol-

water brine solutions at Tc > -40°F,

for Tout ^ -30°F; and Freon-12 or

or other direct expansion coolants

at TC<-40°F, for TQUt < -30°F.C

(12 x rH x p ) * fj.
                                                    Calculate using appropriate equations
                                                    for forced connection in pipes.6
Coolant flow (Wc), lb/hrf

Temperature change of coolant
 (ATC),°F
Coolant flow (Vc),
Clean overall heat transfer
 coefficient (Ur), Btu/ft^-hr-°Fh
757.9 x p

Qtot * (Cp x Wc)


 94.5

[(1.149 * h0) + (0.0000839)
              + (1 * hj]-1
                                       E-45
-------
                                   1   1    •;:  •  ..;•."• f?l;V1"!:f ••'•I	"v	'•'•:';:>'•'if Mr

                  Table E-14.  PROCEDURES TO CALCULATE HEAT TRANSFER
                  AREA OF A CONDENSATION SYSTEM OF STYRENE IN AIR (Continued)
Dirty overall heat transfer
coefficient  (Ud), Btu/ft*-hr-°F

Log nvean  temperature difference  (LMTD), °F


Required  heat transfer area  (A), ft2
 Total  tube  length  required  (Lt),  ftj

 Required heat exchanger length  (LH.E.)>  ft

 Required refrigeration capacity (RC1), tons

 Selected refrigeration capacity (RC), tons
 [(1 * Uc)  + 0.001]

      -AT2) * In
                   -1
Qt t * (UD x LMTD);       :
if A > 43.8 ft2, try a larger
 heat exchanger1
 A * 0.2618

 Lt * 17

 Qtot * 12>000

RC' or minimum of 1
^Condenser and tube characteristics from pp.  11-1  thru 11-18 of  the
 Chemical Engineers' Handbook (Reference 30):
 Tube: outer diameter, D0 = 1.00 in.;  inner diameter, D-j  = 0.870 in.;
         thickness, Xw = 0.065 in.; specific external surface area =
         0.2618 ft2/ft;
         cross-sectional area = 0.004128 fWtube

   Condenser:  shell inside area, Ai = 0.3553 ft*-;
               total tube area, A  = 0.09272 ft*;
               net area = 0.2626 ft2, wetted perimeter = 6.54 ft;
               hydraulic radius, r^ = 0.04001 ft.,
               length, L = 8 ft, total cross-sectional area inside of
               tubes  = 0.004128 ft2/tube x  17 tubes = 0.07018 ft^.

Assuming baffle cuts, lr = 0.25   (shell diameter, Ds); shell outer tube
 limit,  D0tl «7.634  in.   (7/16" clearance for fixed tube sheet for
 Ds < 24");  baffle  spacing, bs = Ds % 8  in.

cCoolant characteristics can  be  interpolated or extrapolated for the
 coolant temperature, Tc,  from   The Chemical Engineers' Handbook:
 pp.  3-71,  206, 213,  &  214 (Reference 34)  for water;  pp.  12-46 thru 12-48
  (Reference 31) for ethylene  glycol water  solutions;  and  pp. 3-191 and
 3-212  thru 3-214  (Reference  31, plus p. E-26  (Reference  32) of The Hand-
 book of Chemistry and  Physics  for Freon-12  (dichlorodifluoromethane).
 Characteristics  required  are dynamic viscosity  (fx),  density (p),
 specific heat (CB),  thermal  conductivity  (k), and specific  gravity (y)
 * p/62.42, lb/fts.

 dFor coolant velocity,  V = 3  fps (3-10  fps recommended by Kern  in Process
  Heat Transfer (Reference  35).
                                         E-46
-------
FOOTNOTES FOR Table E-14 (concluded)
eFrom The Chemical Engineers' Handbook, pp. 10-12 thru 10-15 (Reference 29).
   (1) For turbulent flow (NRe > 10,000) (from Eq. 10-51):
       h  =
              0.023 x V.  ft/hr x  plb/ft3  x  Cn,  Btu-lb-°F   Y
             —   —-                           p_	   x
                        2/3      0 2
                   (NPr)     (NRe)
                                         (\0 44
                                      /id \ *   «1,  if  properties

                                      "V
      at average of bulk  (b)  & wall  (w)  temperature.

  (2)  For transition flow (2000 < NRS  <  10,000)  (from  Eq. 10-49):
       h  _  0.029 k  (NRe2/3 -
        °      rH

   (3) For laminar flow (NRS < 2100) (from Eq. 10-40):
    = 0.465k
                H
where
                      NGz1/3
                       x Npr x 4 x
                                        + 0.87 (1 + 0.015 NQz1/3)
                                        * L.
     coolant velocity of 3 fps and total tube cross-sectional  area of
 0.07018 ft2:  0.07018 ft2 x 180 ft/min. x p , lb/ft3 x 60 min/hr.

9For coolant velocity of 3 fps, 0.07018 ft2 x 180 ft/min. x 7.48 gal/ft3.

hDQ/Di = 1.000 in. 4 0.870 in. = 1.149;
DO Xw =
KtDj_

   where:
              0.0542 ft x 1.000 in.
            69.2 Btu/ft-hr-°F x 0.9335 in.

            Kt = thermal conductivity of brass tube
                 (pp. 23-49 ChE Hndbk) (Reference 26)
      = (D0 - Dj) * In
                                = (l.OO - 0.87) * In (1.00/0.87) = 0.9335.
""See heat exchanger configurations for 1-in. o.d., 1-1/4-in. square pitch,
 T.E.M.A. P or S on p. 11-15 of The Chemical Engineers' Handbook (Reference 30)
 for 8-ft heat exchangers assumed by Enviroscience for cost basis:   17 tube
 minimum unit, A = 35.6 ft2; for 30 tube next larger unit, A = 62.8 ft2;
 assume need larger than minimum size for design (Enviroscience costing
 curve is continuous for all areas) when A = 35.6 + (0.2 heat load  design safety
 margin + 0.1 allowable undersizing) x (62.8 - 35.6) = 43.8 ft2.

JA,ft2 * 0.2618 ft2 of tube external surface area/ft of tube.

kLt,ft of tubes * 17 tubes.

112,000 Btu/hr per ton of refrigeration capacity.

                                  E-47
-------
Table  E-15.   CAPITAL AND ANNUAL OPERATING COST  ESTIMATES
             FOR A  RETROFIT 20 ft2   CONDENSER  SYSTEM  FOR THE
           STREAMS  FROM THE CONTINUOUS  POLYSTYRENE MODEL PLANT
        Item
                                                      Value
      Control  system



      Capital  Cost:

      Purchase cost

      Installed capital cost3

      Annualized cost;

      Operating labor

      Maintenance0

      Utilities:
        Water0

        Electricity6

      Taxes, insurance-and
        administration

      Capital recovery9

      Total annualized cost
        without recovery  credit

      Total amount of styrene
        recovered from W  Ib/hr
        of styrene
       Annual  styrene recovery credit
         at $0.3575/1b

       .Total  annualized cost after credit  ($1,980 - $Z)
Heat exchanger with a maximum
capacity of 20 ft* heat transfer
area
$2,000

 2,960



$1,080

   150


$    5

$  140


$  120

$  480


$1,980
 (W Ib/hr x 8,000 hr/yr x X heat exchanger
 efficiency x 90% recovery efficiency
 from the separator)
 * 2,000 Ib/ton = Y tons/year

 Y tons x 2,000 Ib/ton x $0.3575/1b = $Z
       Cost effectiveness of emission
         reduction  ($/Mg)
 ($1,980 -  $Z)/[W Ib/hr x 8,000 hr/yr
 x X heat exchanger (VOC reduction)
 efficiency/2,205 Ib/Mg]
       Purchase cost times retrofit installation cost factor of  1.48  (see Table 5-2).

       Operating labor cost = 1 hr/wk x  52 wk/yr x 1.15 (with supervision/
        without supervision) x $18/hr (including overtime).

       Maintenance cost = 0.05 x (installed capital cost).
       dWater cost = 25 gpm x 60 min/hr x 8,600 hr/yr x 0.001 make-up/total
        x $0.30/(1,000 gal) x (1,000 gal)/l,000 gal.

       eElectricity consumption (equations from Reference 40) and cost:
        hydraulic horsepower = 50 ft x (1.0 specific gravity) x 25 gpm/3960.» 0.3157 hp

        brake horsepower = 0.3157 hp x 745.7 W/hp x 8,000 hr/yr

                          x kW/l,OOOW +  0.65  pump efficiency  = 2,900 kWh/yr

        Cost - 2,900 kWh/yr x $0.049/kWh


       fTaxes insurance, and administration cost »  0.04 x  (installed capital
        cost).
                                                       ;     '  "iGi'»  M   '         ''
       9Capital recovery factor = 0.1627, for ,10 percent interest (before  taxes)
        and 10 year life.

                                       E-48
-------
          Table E-16.   CAPITAL AND ANNUAL OPERATING COST ESTIMATION
                        PROCEDURE FOR CONDENSERS WITH REFRIGERATION
           Item
       Value
Capital Costs
    Condenser
      Installed cost, Dec. 1979

      Installed cost, June 1980a

      Installed retrofit cost, June 1980


    Refrigeration
      Installed cost, Dec, 1979

      Installed cost, June 1980a

    Total Installed Cost, June 1980

Annualized Costsc
    Operating labor^

    Maintenance materials & labor

    Utilities
      Electricity, pumping

      Electricity, refrigeration

      Coolant, make-up

    Capital recovery!1

    Taxes, administration
      & insurance

    Total annualized cost
      without recovery credit
    Styrene recovery credit

    Net Annualized Cost
       after recovery credit
From Figure E-6 for A

Installed cost, Dec. 1979 x 1.047

Installed cost, June 1980 x 1.065,
from Table 5-2


From Fig. E-7 for RC & Tcb

Installed cost, Dec. 1979 x 1.047

Sum of condenser and refrigeration


$1,080

0.05 x total installed cost


See footnote e

See footnote f

See footnote g

0.1627 x total installed cost

0.04 x total installed cost
Operating labor & maintenance
  + utilities + capital recovery
  + taxes, administration &
    insurance

2767 x Ws x (% Red'n. * 100)

Total annualized cost - styrene
  recovery credit	
                                    E-49
-------
Footnotes for Table E-16
                                                  ll'i'i,:1
aUpdated using Chemical Engineering Plant Cost  index  from  December 1979
 (247.6) to June 1980  (259.2).
                              . j.    ,   .. ••'  	''  ,.', I&v'ih  -\ .  '!:j •.:: ••;<'•:.
"Costs for the 1 ton minimum  refrigeration  capacity can be approximated
 by exp (exp[(0.60784  x In  (hp/ton)) +  0.31169])  .,
           t: '               '  ' !  '  '• •:	'  '  •'   ;•	 -:*w	[:y.j:  •'':	:	
cCost factors presented in  Table  5-3.

dOperating labor cost  = 1 hr/wk x 52 wk/yr  x  1.15  (with supervision/without
 supervision) x $18/hr (including overtime).

eUsing Equation 6-2, p. 6-3 in The Chemical Engineers'  Handbook
 (Reference 40) for V  = 3 fps (for condensers with a  heat  transfer area
 of 20 ft2 or less and 125  ft2) or 10  fps  (for  condensers  with  a heat
 transfer area of 185  ft2), assuming  a  pumping  height of  50 ft. and a pump
 efficiency of 65%:
                                                                             	i", .  II-. !!il '•':
          50  ft  x V x  Vc
          3960 gpm  ft/hp
x 0.7457 kW  x
      Fp
    8000 hr/yr
0.65 pump efficiency

   x $0.049/kwh
           " ,                    i        i               i'  i
           where 7= specific gravity of coolant = p* 62.42 lb/ft3
                                                   (the density of water)
                             •  I             '    •  I;IM si'[?,' •"]•,'' ""]""'	''i"  '","'' •  ."'
                Vc = volumetric flow of coolant; equal's 94.5 gpm for
                     condensers with heat transfer area of 20 ft* OP
                     less;  equals 472 gpm for condensers with heat
                     transfer area of 125 ft2; and equals 1,575 gpm for
                     condensers with heat transfer area of 185 ft?

 f        RC'  x (hp/ton of refrigeration for TC)	
      0.85 compressor efficiency x 0.85 motor eTficiency

         x 0.7457 KW  x 8000 hr/yr x $0.049/kwh
              hp

         where (hp/ton of refrigeration) for a particular coolant temperature
         is given on Fig E-7 for multiples  of 20°F between -60 and + 40°F
         or can be calculated from the curve fit:
                                                                      i
          (hp/ton) =  exp [-0.1777 + 0.01503 (45-T)]

 9For chilled water, assume 99.9% recycle:
                    @ 94.5 gpm x 60 min/hr  x 8000 hr/yr x 0.001 make-up
                                            x $0.30/1000 gal  =  $14/yr, ,
                                             use $20/yr;
                                    E-50
-------
Footnotes for Table E-16 (Concluded)

 For ethylene glycol-water brine solutions and Freon-12,  assume one
 replacement per year of coolant in condenser and refrigeration system
 and coolant volume in condenser and refrigeration twice  that  of condenser
 alone.

       Coolant volume, gal = A x 0.004128 ft2 x-sect./tube
                           x 7.48 gal /ft3 x (2 x inside
                           tube volume in condenser) *
                           0.2618 ft2 surf ace/ ft tube
                           = 0.2359 x A

For ethylene glycol-water brine solutions:

       cost of coolant = Xw ($0.30/1000 gal) + X£G  ($0.27/lb xpEG lb/ft3

                         * 7.48 gal /ft3)

                       = $0.0003 XN + $2.02
       where:  X   = volume fraction of water in brine solution,
                   = volume fraction of ethylene glycol in brine solution
 For  Freon-i2  solutions:
         cost of  coolant = $8.70/liter x 3.785 liter/gal based on 20 liter
         lot price  of  trichlorotrifluoroethane reagent price of $8. 73/1 Her
         from Fisher Scientific Co. 1979.

 n!0 percent interest  (before taxes) and 10 yr. life.

 ""W.,  Ib styrene  emitted/hr  x 8000 hr/yr x (% Red'n  in condenser * 100)
         x 0.90,  fraction of reduction recovered x $0. 3575/1 b styrene.
                                   E-51
-------
  io,ooo
   1,000
03
•»J
•I—
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•o
-------
  1,000
o
o
o
(O
a.
to
O
(U
(O
-p
to
C
i.
CD
-Q

OJ
O
O)
Q
100 -
                               10                      10O

                 Refrigeration Capacity - Tons  (12,000  Btu/hr)


          Figure E-7.   Installed Capital Costs  vs.  Refrigeration
               Capacity at Various Coolant Temperatures for a
                        Complete Refrigeration  Section
                                                                          1,000
                                    Er53
-------
based on an equation in the Chemical  Engineer's  Handbook41  and  simplified
as suggested by Chontos.42»43>44  The next larger  size (inner diameter)
of schedule 40 pipe was selected unless the calculated size was within
10 percent of the difference between  the next smaller and next  larger
standard size.  For flows of 700 scfm and greater,, duct sizes were
calculated assuming a velocity of 2,000 fpm for flows of 60,000 acfm
or less and 5,000 fpm for flows greater than 60,000 acfm.  Duct sizes
that were multiples of 3-inches were used.
E.6.2  Piping and Ducting Cost Estimation Procedure
     Piping costs were based on those given in the Richardson  Engineering
Services Rapid Construction Estimating Cost System18 as combined for
70 ft. source legs and 500 ft. and 2,000 ft. pipelines for the  cost
analysis of the Distillation NSPS.45  (see Tables E-18 and E-19)
Ducting costs were calculated based  on the installed cost equations
given  in the 6ARD Manual.46  (See Table E-20.)
     Costs  of  source  legs were taken or  calculated directly from the
tables.  Costs of pipelines  for  flares were  intefpotated for the safe
pipeline  lengths differing  by more than  10 percent from the standard
lengths of  70, 500, and  2,000 ft.  Installed capital  costs were updated
to June 1980  using  the Chemical  Engineering  pipes'!! valves, and  fittings
index  for piping and  the fabricated  equipment  index  forducting.
                                   E-54
-------
                   Table  E-17.   PIPING AND  DUCTING DESIGN PROCEDURE
         Item
                                                                          Value
 (1)  Pipe diameter, D

     (a)  Piping3
     (b)  Ducting13
(2)  Pipe length,  L

     (a)  Flares
     (b)  Incinerators
                                                                       + 0.472, for Q<40 scfm
                                                                       + 2.85, for 40 12 in. or Q>700scfm
 and Q^ 60,000 acfm
 D  (in.) - (0.1915)7Q(acfm),  for D >60,000 acfm
 Select size that is a multiple of 3 inches.
Assumed 70-ft. source leg  from  each  source
to the pipeline.
Assumed separate pipelines for  large  (>40,000 scfm)-"
intermittent streams and for  all continuous
streams together.  Selected pipeline  length of
70, 500 or 2,000 ft. if calculated safe pipeline
length within 10 percent of standard  length; if
not selected calculated length  between
standard values.

Assumed 70-ft. source legs from each  source
to the pipeline.
Used duct, fan. and stack  cost  from
Enviroscience,'° which assumes  a 150-ft.
duct cost based on  the CARD Manual
(Reference 46)
Economic  pipe diameter equations from Reference 44 (which  is based upon References 41
 and 42).

 From continuity equation Q= -| D2V; assumed velocity, V, of 2,000 fpm for lower flows
 and 5,000 fpm for higher flows.
                                             E-55
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E.7  REFERENCES FOR APPENDIX E

 1.  Kalcevic, V.  Control  Device Evaluation:   Flares  and  the Use  of
     Emissions as Fuels.   In:   Organic Chemical  Manufacturing Volume 4:
     Combustion Control  Devices.   U.S. Environmental Protection Agency.
     Research Triangle Park, N.C.  Publication No.  EPA-450/3-80-026.
     December 1980.

 2.  Reference 1, p. IV-4.

 3.  Memo from Sarausa,  A.I., Energy and Environmental  Analysis, Inc.
     (EEA), to Polymers  and Resins File.  May  12,  1982.  Flare costing
     program (FLACOS).

 4.  Telecon.  Siebert,  Paul, PES with Straitz, John III,  National Air
     Oil Burner Company, Inc. (NAO).  November 1982.   Design, operating
     requirements, and costs of elevated flares.

 5.  Telecon.  Siebert,  Paul, PES with Fowler, Ed,  NAO.  November  5,
     1982.  Purchase costs and operating requirements  of elevated
     flares.

 6.  Telecon.  Siebert,  Paul, PES, with Keller, Mike,  John Zink Co.
     August 13, 1982.  Clarification of comments on draft  polymers and
     resins CTG document.

 7.  Telecon.  Siebert,  Paul, PES with Fowler, Ed,  NAO.  November  17,
     1982.  Purchase costs and operating requirements  of elevated
     flares.

 8.  Memo from Senyk, David, EEA, to EB/S Files.  September 17, 1981.
     Piping and compressor cost and annualized cost parameters used  in
     the determination of compliance costs for the EB/S industry.

 9.  Memo from Mascone, D.C., EPA, to Farmer,  J.R., EPA.  June 11,
     1980.  Thermal incinerator performance for NSPS.

10.  Blackburn, J.W.  Control Device Evaluation:  Thermal  Oxidation.
     In:  Chemical Manufacturing Volume 4:  Combustion Control Devices.
     U.S. Environmental  Protection Agency, Research Triangle Park,
     N.C.  Publication No. EPA-450/3-80-026, December  1980.
     Fig. III-2, p. III-8.

11.  Reference 10, Fig. A-l, p. A-3.

12.  Air Oxidation Processes in Synthetic Organic Chemical Manufacturing
     Industry - Background  Information for Proposed Standards.   U.S.
     Environmental Protection Agency, Research Triangle Park,  N.C.
     Draft EIS.  August 1981.  p. 8-9.

13.  EEA.  Distillation NSPS Thermal  Incinerator Costing Computer
     Program  (DSINCIN).  May 1981.  p. 4.
 14.  Reference 10, p.  1-2.
                                  E-59
-------
                                                                          i	,.,'jm.: jiii-'
15.  Reference 12, p. G-3 and 6-4.

16.  Reference 10, Fig. V-15, curve 3, p. V-18.

17.  Never-ill, R.B.  Capital and Operating Costs of Selected Air
     Pollution Control Systems.  U.S. Environmental Protection Agency,
     Research Triangle Park, N.C.  Publication  No. EPA-450/5-80-002.
     December 1978.

18.  Richardson Engineering  Services.  Process  Plant  Construction Cos
     Estimating Standards, 1980-1981.  1980.

19.  Telecon.  Katari, Vishnu,  Pacific Environmental  Services,  Inc.
     with Tucker, Larry,  Met-Pro  Systems  Division.  October 19,  1982.
     Catalytic incinerator  system cost estimates.

20.  Telecon.  Katari, Vishnu,  Pacific Environmental  Services,  Inc.,
     with Kroehling, John,  DuPont, Torvex Catalytic Reactor Company.
     October  19,  1982.   Catalytic incinerator system  cost estimates
         "•'"     .     .       ,  i	     . ••   I- / ,  '	 . :<[,;,' ,!	i1 •	j,i",s" '••< - i
                              I      " , ,  ,•'.'..•.',• '.i , ••.':	, idlii	K..'1;1 j'i  V •
21.  Letter  from  Kroehling,  John, DuPont, Torvex Catalytic Reactor
     Company,  to  Katari, V., PES.  October 19, 1982.   Catalytic incinerator
     system  cost  estimates.

22.  Key, J.A.  Control  Device Evaluation:  Catalytic Oxidation.  In:
     Chemical  Manufacturing Volume 4:   Combustion Control Devices.
     U.S.  Environmental  Protection Agency.  Research Triangle Park,
     N.C.  Publication No.  EPA-450/3-80-026.  December 1980.

 23.  Telecon.  Siebert, Paul, Pacific EnvironmentalServices, Inc.,     _'
     With Kenson, Robert, Met-Pro Corporation, Systems Division.  July  22,
      1983.   Minimum size catalytic incinerator units.
           '   •                li  . . j f: ,  '  '   " '     ",  ii'niili i ','	l,''l,   , 'i < i,  ! HI,	 ' 	  '  »,, ', I' -I  i,1 'i '<
                              ^        '        .   '.ILj-filit.  i ,"  -.' ::;'.". 	v:.. .'.:,• 	 t1,,, -i:..'^ 'i'-l
 24.   Steam:   Its Generation and Use.  New York, Babcock  &  Wilcox Company,
      1975.   p. 6-10.

 25.   Reference 10,  Fig.  V-15,  curve 3, p. V-18.

 26.  Perry, R.H. and C.H. Chilton, eds.   Chemical Engineers' Handbook,
      fifth edition.  New York, McGraw-Hill  Book  Company.  1973.   p.  3-59

 27.  Rossini, F.D.  et al. Selected Values of Physical  and  Thermodynamic
      Properties  of  Hydrocarbons  and Related Compounds,  Comprising the
      Tables of API  Research Project 44.   Pittsburg,  Carnegie Press,
      1953.   pp.  652 and 682.

 28.  Reference 26,  pp.  10-25 through  10-28.

 29.  Reference 26,  pp.  10-12 through  10-15.

 30.  Reference 26,  pp.   11-1 through  11-18.
.ill I!	":"!!: "I",III I
    I
                                                                        •	M1! "i-a.
                                    E-60
-------
31.  Reference 26,  pp.  3-191,  3-212  through  3-214, and 12-46 through
     12-48.

32.  Weast,  R.C., ed. Handbook of Chemistry  and  Physics,  fifty-third
     edition.  Cleveland, The  Chemical  Rubber  Company.  1972.   p. E-26.

33.  Erikson, D.G.   Control  Device Evaluation: Condensation.   In:
     Organic Chemical Manufacturing  Volume 5:  Adsorpiton,  Condensation and
     Absorption Devices.  U.S. Environmental Protection Agency, Research
     Triangle Park, N.C. Publication No.  EPA-450/3-80-027. December  1980.
     p. A-3.

34.  Reference 26,  pp. 3-71, 3-206,  3-213, and 3-214.

35.  Kern, D.Q.  Process Heat Transfer.  New York, McGraw-Hill  Book
     Company, 1950.  p. 306.

36.  Telecon.  Katari9 Vishnu, Pacific Environmental  Services,  Inc.,
     with Mr. Ruck, Graham Company.   September 29,  1982.   Heat  exchanger
     system cost estimates.

37.  Telecon.  Katari, Vishnu, Pacific Environmental  Services,  Inc.,
     with Glower, Dove, Adams Brothers, a representative  of Graham
     Company.  September 30, 1982.  Heat exchanger system cost estimates.

38.  Telecon.  Katari, Vishnu, Pacific Environmental  Services, Inc.,
     with Mahan, Randy,  Brown Fintube Company.  October 7, 1982.  Heat
     exchanger system cost estimates.

39.  Reference 33, pp. A-4 and A-5.

40.  Reference 26, p. 6-3.

41.  Reference 26, p. 5-31.

42.  Chontos, L.W.   Find Economic Pipe Diameter via Improved Formula.
     Chemical Engineering.  87(12):139-142.  June 16, 1980.

43.  Memo from Desai, Tarun,  EEA, to EB/S Files.  March 16,' 1982.
     Procedure to  estimate piping costs.

44.  Memo from Kawecki,  Tom,  EEA, to SOCMI Distillation File.   November. 13,
     1981.   Distillation pipeline costing model  documentation.

45.  EEA.   SOCMI Distillation NSPS  Pipeline Costing Computer Program
      (DMPIPE), 1981.

46.  Reference 17,  Section  4.2,  p.  4-15  through 4-28.
                                   E-61
-------
-------
                APPENDIX F
CALCULATION OF UNCONTROLLED EMISSION RATES
     AT SPECIFIC COST EFFECTIVENESSES
                   F-l
-------
APPENDIX F.
                        CALCULATION OF UNCONTROLLED EMISSIONRATES
                        AT SPECIFIC COST EFFECTIVENESSES
                               • :    '      ;   (.;..,,  '   ' ill I   i  : -,  - .  , ,.	,,  . i  i
     This appendix details the procedures used to calculate the  uncontrolled
emission rates equivalent to $1,000 per Mg, $2,000 per Mg, and $3,000 per Mg
when RACT is applied.  Section F.I describes the procedures for  flares,
thermal incinerators, and catalytic incinerators.  Section F.2 describes the
procedures for condensers.                                             ;
F.I.  PROCEDURE FOR  INCINERATION DEVICES
     For the polypropylene and high-density polyethylene model process
sections, the question asked was what  uncontrolled VOC emission  rates when
reduced 98 percent  (i.e., RACT level)  corresponded to cost effectivenesses  of
$1,000 per Mg, $2,000 per Mg, and $3,000  per Mg.  fne foil owing  sections
describe the procedures used to calculate these  uncontrolled  emission rates.
F.I.I  General Procedure
     The general  procedure used is  as  follows:
     First.  For  each process section  identified in  Tables 4-1 and  4-2,
the emission characteristics  identified  in  Chapter 2 of  the CTG  and Chapter 6
of  the background information document for  the  polymer manufacturing industry
            J                    ,                  	     ...         i   	
and the  control  costs identified  in Chapter 5  of the CTG were used  as  the
starting point.   Table  F-l summarizes  the pertinent  information.
     Second.   Uncontrolled emissions were adjustedproportionally by changing
volumetric  flow  proportional  to the initial flow.  Concentration of the
emissions was  assumed to  remain  constant.  Uncontrolled  emissions needed to
be  adjusted downward or upward  depending upon the  initial  cost  effectiveness.
For example,  if  the initial  cost effectiveness was  $l,500/Mg, the uncontrolled
 emissions would  be higher than  the initial  uncontrolled  emissions in order to
                                                                       j
 correspond  to $l,000/Mg and  lower than the initial  uncontrolled emissions in
 order to correspond to  $2,000/Mg and $3,000/Mg.
      Third.  Annual costs were adjusted to take into account the new flow
                                i '         '  ..'>.( ,.•'  '.[Jir:1 ;..•!; '« , I',!1," •'•!  •' 'I   "' '••)	II' I". ."i:V Ml!';;
 conditions, which affect control  device costs.  Annual  costs were divided
 into three  components:  (1) those related to'"capital  costs (C^,  (2) those
 related to operating costs (C2J, and  (3) minimum and/or constant costs
                     F-2
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                                                                               I
                                                                      ;'!	^iORI'
Table F-2 summarizes these costs for each of the two polymers.  The annual
costs were adjusted as follows:
     o  Capital-related costs, C]_, were adjusted using the equation
            i0'6 where: Vj = total initial volume of flow from model process
                            section and
                      Vx = adjusted flow rate.
     o   Operating-related costs, C2, were adjusted using the equation

          ^  '             ',                 .  '.   ,   '. '   ', '        „,  .,'.2',
     Capital-related costs included capital  recovery", maintenance,
                           1           ,    ,;	 '  , ,,!«„ "», ' !  'I1  r , i ! • • ,  "  „ ' .'  | "   „• "« : " ,•!! -'hi 	I
taxes, insurance and administration charges.  Operating-related costs
included utilities  (e.g., natural gas, steam, electricity).   Operating
labor was assumed to be constant.                                 \
     Fourth.  As flow rates vary, the  size of  the control device required
will also vary.  No matter how  small  the flow, however, there are
certain minimum size control  devices  available; thus, control device
costs do not  approach  zero as flows become  very small.  In addition,
some utility  requirements, such as  natural  gas  purge  rates,  may be
constant, or  even increase as flow  rates become increasingly smaller.
Finally, a  different  control  device design  may  be  more  cost  effective
as  flow  rates change.   For example, as  flow rates  approach 1.46  sefm
(70°F),  a change in flare  design was  assumed to occur where  a flare
with a  fluidic seal  was used  for flows  less than 1.46  scfm.   Table F-3
summarizes  the basic minimum  costs  associated with the  various  control
devices  at  flow rates  that affect  design criteria.               i
     Fifth.  Using  the above  information and procedures,  the following
basic  equation was  solved  for Vx for  each  process  section:
         c.  /Vx\0-6  +  C^'/Vx',.   ~,
                                     3 = $l,000/Mg; $2,000/Mg;
                                    —   and $3,000/Mg
 where:
C1  = the difference between the capital-related costs of
 1   the control device controlling V-^ and the capital-related
     costs of the control device controlling V2-
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Table

Control
Device
Flare, within
line



Flare, across
lines





Thermal
Incinerator
- within
line
- across
lines

Catalytic
Incinerator
- within
line


- across
lines



F-3.

Flow,
scfm

1.49
1.46
0.81
0.03


1.49
1.46
0.81
0.03




32.19

96.58



754.6

150

500

150


BASIC MINIMUM COSTS
',. ' .';,. '.;

Capital -Related

4,887b
4,964d
4,964d
4,964d


6,208b
d,304d
6,304d
6,304d
1,



75,240

78,000
i'"


42,260

37,260

45,480

43,060
1 i : ,',
t . ' ' , 1; ! '- .'
if, ', ,• i
AT VARIOUS FLOW
', ANNUAL CSSJS, $

Operating-Related
in i i :
4,276C
4,276C
4,276C
4,276C


4,276C
4,276C
4,276C
4,276C
: • . " l! " '. ' ! ' • "'



55

160 ;:
,, , .. , 	
	


3,940

	 i,470'i:'""
'', , , v i " 1 ' ' ,1
2,620

1,730'
• ""• ' ' , „ ' ' i1"'

RATES3
- '- 	

Constant

11,160
11,160
11,160
11,160


11,160
11,160
11,160
11,160




21,600
11 " I!'1" "'
21,600




16,700

16,700
1 «" 	 ,
16,700

16,760'



11 	 !| 	 I1! 	 /Iji'.'!'!^!!!:!^'!^1!!1!!!''! 	 Hill
,' ' "•••! 	 f«f ' -ii
1 III



Total
•'' I
'•.f,11
26.323C
26,400C
20,400C
20,400C


2^,644°
2i,740c
2l,740c
21.74QC
'i ' 1




96,895

	 ''' |M !"
99,760


- ' 	 ; ' JJ ' i1
	


62,900

'55,430 	 '

64,800

61,490
. '••
                                                                                                      	I'liejiis:
                                                                                                      -	f	''	
                                                                                                        i	a
                                      ,1  .   " i            • ••<  ', '  .. M'tl ',. ••!•  , I     .""'II i,          I
 The minimum costs for flares are based on a single emission stream  (i.e., one source  leg); per
 process section.  If more than one emission stream emanates from a  process  section, then 'minimum
 capital costs will be higher than those reported in the table.  The increase in  capital-related
 costs is about $690 per additional source leg at 1.49 scfm and about $670 per additional^source
 leg at lower flows.  The minimum incinerator costs are specific to  the  process sections for
 which they were costed.                                                                  >

 Flare without a fluidic seal.       '                                                     '       '

CAdd steam costs at 1.49, 1.46, 0.81, and 0.03 scfm.  Actual cost is dependent on molecular weight
 of gas stream and weight percent of VOC.
d
 Flare with fluidic seal.
                                            F-6
-------
        C2
        c1
         3
= the difference between the operating-related costs for
  the control  device controlling V-^ and the operating-
  related costs of the control  device controlling V2«
= the minimum costs associated with controlling V£.
        C'  =  the difference  in  emission  reduction  associated with
         4    controlling  emissions  at  V-^ and  emissions  at  V£.
        C'  =  the emission reduction at V?.
         5

        V  =  the initial, or higher, flow rate.

        V  =  the flow rate at the lower  end of the design  range.

        Vx  =  Flow rate to be solved for.
     Table  F-4 summarizes the coefficients used in the  calculations.
     Sixth. Once the flow rates were found, the uncontrolled  emission
rates were  calculated by  the following equation:
                     Vx  x ER
     where:
          Vx = flow rate at $l,000/Mg ($2,000/Mg., $3,000/Mg)
          Ml = initial flow rate from process section
          ER = initial uncontrolled emission rate.
     Table F-5 summarizes these results.
F.2.  PROCEDURES FOR CONDENSERS
     In calculating the uncontrolled emission rates for polystyrene,
the general question that was asked was:  What uncontrolled emission rate,
when controlled to 0.12 kg VOC per Mg of product (i.e., to the RACT level),
yields a cost effectiveness of $1,000 per Mg ($2,000 per Mg, and $3,000
per Mg)?  This is slightly different from incineration where, regardless
of the uncontrolled emission rate, 98 percent VOC reduction was assumed.
For polystyrene, the effective percent emission reduction varies as the
uncontrolled emission varies.  The following paragraphs detail the
procedures used to calculate the uncontrolled emission rates associated
with the three cost effectivenesses.
                               F-7
-------
Table F-4.  SUMMARY OF COEFFlCtelts
Cost
Process Effectiveness, Coefficient
Polymer Section S/Hg Vj. V2 q ^2 C3
Polypropylene 1.00°
- within line RHP 2,000
3,000
1,000
PR 2,000
3,000
1 ,000
MR 2,000
3,000
0.81 0.0627 0 212 20,418

0.81 0.03 0 114 20,404

0.81 0.03 0 138 22,415

\ '" 1 000
PF 2,000
3,000
- across lines 1.00°
886.3 32.19 20,160 1,455 96,895

0.81 0.188 0 177 21,793
RHP slooo 0.188 0.03 0 44 21,749
1,000 1.46 0.81 0 95 21,858
PR 2,000
3,000
	 • 	 i ,000
MR 2,000
3,000
1,000
PF 2,000
3,000
0.81 0.03 0 114 21,744
i, ' , "'. ,''",.' i' i , "' ,
1.46 6.81 0 IT3 27,916
0.81 0.03 0 138 27,778
! • "„ ::,; ";;,: ; '
2658.8 96.58 26,640 4,370 99,760

1 ; i "" • "' •.<" '. « ;;, .'"4 .if i 	 11! IS1, • 	 I ;" :,,"
High-Density
Poly^htylene 	
-within line 1,000 69-63 l.« 2,349 6,099 20,456
m 2>00° | 0.81 0.03 0 69 20,403
3,000 >
[ j 	 " i ;' 	 ! 	 , ':, "'I1,!1'. 	 i'1 	 "'! ' " ' "
1,000 1754.6 251.5 4,370 2,390 '' 56 1 140'
PF 2>°°° I 251.5 150 630 80 55,430
3,000 )
1,000 208.9 1.49 7,555 18,565 ' 21,834
MR 2,000 1.46 0.81 0 59 21,812
3,000 0.81 0.03 0 69 21,743
- across lines 	
1,000 754.6 500 5,480 1,320 64,800
PF 2 ,000 ]
(500 150 2,420 890 61,490
3,000 ) , : ;, .

	 c\ 	 .,;.
38.428

20.376

24.796

• • •
115.406

32
8.129
16.985
20.376

24.796

346.218

.;; „ ' | i 	 ',

868.41
9.94

i:i' hi ,i r.iii , ,:!",
56.738
11.449

2,644.45
8.29
9.94
28.728
39.493


°f 	
3.224

0.I784

0.:954


4J35

9*67
1,5435
21,16
01784
'i
0.954

13.05



1EJ.99
0.38


281 .369
15.92

18.99
10.32
0.38
56.418
16.925


               F-8
-------
Table F-5.  SUMMARY OF COST EFFECTIVE FLOWS AND EMISSION RATES,
       POLYPROPYLENE AND HIGH-DENSITY POLYETHYLENE PLANTS
Polymer
Polypropylene
- within line
- across lines
High-Density
Polyethylene
- within line
- across lines
Process
Section
RHP
PR
MR
PF
RMP
PR
MR
PF

MR
PF
MR
PF
Cost
Effectiveness
1,000
2,000
3,000
1,000
2,000
3,000
1,000
2,000
3,000
1,000
2,000
3,000
1,000
2,000
3,000
1,000
2,000
3,000
1,000
2,000
3,000
1,000
2,000
3,000

1,000
2,000
3,000
1,000
2,000
3,000
1,000
2,000
3,000
1,000
2,000
3,000
Vx * VX
0.3989
0.1989 0.0627
0.1325
0.7852
0.3915 7.175
0.2608
0.7088
0.3535 43.46
0.2354
875.85
406.39 886.3
262.67
0.4249
0.2119 0.188
0.1411
0.83665
0.4172 21.526
0.2779
0.8785
0.438 130.4
0.29175
841.6
398.4 2,658.8
259.35

1.6095
0.8033 69.63
0.535
537.4
248.8 251.5
164.38
1.7245
0.856 208.9
0.5701
608.4
279.9 754.6
133.67
Uncontrol led
Emission
x ER = Rate
0.445
0.07 0.222
0.1479
0.445
4.1 0.222
0.1479
0.489
30 0.244
Ovl62
2.569
2.6 1.192
0.771
0.1582
0.07 0.0789 .
0.0525
0.1582
4.1 0.789
0.0525
0.2021
30 0.1007
0.0671
0.823
2.6 0.389
0.254

0.2936
12.7 0.1465
0.0976
0.8675
0.406 0.4016
0.2654
0.1048
12.7 0.052
0.0347
0.3273
0.406 0.1506
0.0988
Annual
Emissions
(x relevant
capacity)
20.93
10.44
6.95
20.93
10.44
6.95
22.996
11.468
7.637
120.74
56.02
36.237
22.31
11.12
7.41
22.31
11.12
7.41
28.497
14.21
9.464
116.04
54.849
35.814

20.93
10.45
6.96
61.85
28.63
18.92
22.44
11.14
7.42
70.04
32.23
21.14
                              F-9
-------
                                             'f Si?:1 'Wi	If	" ITS	'1;:':'';!•
F.2.1  Styrene-in-Steam Emissions
                               .  -    .  ,    ,•. ,  ,  ,       ,:,' , .  ,.      	>,
     The basic equation for calculating cost effectiveness is" as"as follows:

     (1)     CE = AC - (0.9 ERed x RC)
                          ERed"
     where:  CE   = cost effectiveness, $/Mg
             AC
                                                     iil*!.
                  = annual!zed cost of condenser to reduce uncontrolled
                    emissions to 0.12 kg VOC/Mg product,  $/yr
             RC   = recovery credit, $/Mg, = $0.788/kg of styrene     !
             0.9  = efficiency of actually recovering the styrene
                    from the condenser
             ERed = annual emission reduction from uncontrolled
                    to 0.12 kg VOC/Mg product, Mg/yr
For polystyrene, we are already dealing with a minimum-size condenser
and operating requirements (which were assumed constant) when the
uncontrolled emission rate is at 3.09 kg VOC/Mg product.   In order to
get a cost effectiveness of $1,000  per Mg, a smaller uncontrolled
emission rate is needed.  Thus', annualized costs associated with
polystyrene are a constant -  equal  to $8,300.
     Emission reduction,  in  general, can  be  calculated with the
following equation:
     (2)     ERed  =  (Emission Rate x Capacity)  -  (0.12 x Capacity)
Capacity is  given:   36.75 Gg for  a process  line and 73.5 Gg
for the plant.   Thus the above  equation reduces to:
     For single process line:
      (3)     ERed = (ER x 36.75)  - (0.12 x 36.75)
                   = 36.75 ER -  4.41
      For  two process lines:
      (4)      ERed = (ER x 73.5) -(0.12 x 73.5)
                   = 73.5 ER - 8.82
                                F-10
-------
Inserting  the above  information  into  the  general cost-effective equation
(1),  the following equation  is derived:

     For a single process  line:
(5)
                = $8,300 -  [36.75  ER -  4.41]  (0.9)  [$788/kg]
                               (36.75 ER -  4.41)
     For two process lines:

     (6)     CE = $8.300 - [73.5 ER - 8.82]  (0.9)  [$788/kg]
                               (73.5 ER - 8.82)

     As we know CE (i.e., $l,000/Mg; $2,000/Mg;  or $3,000/Mg),  we  can
solve directly for ER.  Simplifying the above equations  (5 and  6),  we
get:
     For a single process line:
     (7)     ER =  11, .428 + 4.41 CE
                   26,063 + 36.75 CE
     For two process lines (i.e., the model  plant):
     (8)     ER = 14.555 + 8.82 CE
                  52,126 + 73.5 CE
     Substituting $l,000/Mg, $2,000/Mg, and  $3,000/Mg into the  last two
equations (7 and 8), yields the following results:
                 Emission Rate, kg VOC/Mg Product

$l,000/Mg
$2,000/Mg
$3,000/Mg
Single Line
0.2521
0.2034
0.1809
Two Line
0.1861
0.1617
0.1504
F.2.2 Styrene-in-Air Emissions
     As with styrene-in-steam emissions, the basic equation for calculating
cost effectiveness is as follows:
                                  F-ll
-------
     (9)
CE   =  AC - (0.9 ERed x RC)
                ERed
where:
CE   =

AC   =

0.9  =

ERed =

RC   =
                     cost effectiveness, $/Mg

                     annualized costs, $/yr

                     efficiency of collecting recovered styrene

                     annual  emission reduction, Mg/yr

                     recovery credit, $/Mg of styrene recovered

     In calculating the cost effectiveness numbers and the uncontrolled
emission rates for the "across line" analysis, costs were initially
developed for two uncontrolled emission rates: 0.2 kg VOC/Mg product
and 0.15 kg VOC/Mg product.  The resulting costs are summarized in

Table F-6.                '     |  '  '   '      '      V^,/  |/"	." '
     As seen in Table F-6, the uncontrolled emission rates associated

with $l,000/Mg, $2,000/Mg, and $3,000/Mg lie between 0.2 and 0.15 kg
VOC/Mg product.  Using the general equation (9) above and assuming that

refrigeration electricity and recovery credit vary proportionally with emission

rate, the following equation is developed:
                                                                           111	', II	II1'! K iHI" ill!1"1 •'
 where:     AC'     = Constant  costs  associated  with an uncontrolled  emission
                    rate of 0.15 kg VOC/Mg product

           RElec  = Refrigeration electricity  associated  with  an  uncontrolled
                    emission  rate of 0.2 kg VOC/Mg product

           RElec1  = Refrigeration electricity  associated  with  an  uncontrolled
                    emission  rate of 0.15 kg VOC/Mg product

           RC


           RC1

           ERed
     = Recovery credit,  $/yr,  associated with an uncontrolled
       emission rate of  0.2 kg VOC/Mg product
     = Recovery credit,  $/yr,  associated with an uncontrolled

     * Annual  emission reduction  associated with an uncontrolled
       emission rate of  0.2 kg VOC/Mg product
           ERed1   = Annual  emission reduction associated with an uncontrolled
                    emission rate of 0.15 kg VOC/Mgproduct

           ER     = Emission rate to be solved for,, kg VOC/Mg product
                                   F-12
-------

























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F-13
-------
     Substituting the values from Table  F-6  into  equation  10  and then
simplifying* yields the following equations:
                               '            ' "j '  • ' ' i' fit "     ' '     ' '
     For across lines:
(11)
                   = 16.065 - 52.000 ER
                       73.5 ER - 8.82
     For within line ($3,000/Mg only):
     (12)
           CE = 12.950 - 26.000 ER
                36.75 ER - 4.41
     Solving for ER, equations 11 and 12 become:
     For across lines:
      (13)
           ER = 16.065 + 8.82 CE
                52,000 + 73.5 Ch
      For within  line  ($3,000/Mg only):
      (14)
           ER = 12,950 + 4.41 CE
                26,000 + 36.75 CE
      For within line uncontrolled  emission  rates equivalent to $l,000/Mg
 and $2,000/Mg,  the calculations  are  complicated by the changing size of
 the condenser.   (All previous  condenser  calculations assumed the use of
 a minimum size  condenser).   A  cost-effectiveness equation was developed
 to calculate the particular uncontrolled emission  rates; which are
 known to lie between 3.09 kg VOC/Mg  product and 0.2  kg VOC/Mg  (see  last
 column in Table F-6).
      To start developing this  equation,  we  know that at  minimum the
 costs and emission reduction are equivalent to those when the  emission
 rate is 0.2 kg VOC/Mg product.  These' numbers provide a  base,  or minimum,
 from which to start.  In addition, we know that  at most  costs  and    i
 emission reduction  are equivalent to those when  the emission rate  is;
 3.09 kg VOC/Mg product.  The primary calculation is to determine  how
 costs vary  from the minimum (i.e., at 0.2 kg/Mg)  to the maximum (i.e.,
 at  3.09  kg/Mg).   We know that emission  reduction,  and, thus, recovery
 credit,  is  proportional to the emission factors;  that is, there is a
 linear  relationship between emission factor, emission reduction,  and;
 recovery credit.   However, the annualized costs associated with increasing
                                    F-14
-------
 condenser sizes  are not necessarily linear.  Therefore, we assumed that
 the  costs varied exponentially with the ratio of the emission factors
 between  the costs at 3.09 kg/Mg and 0.2 kg/Mg.  The exponents were
 calculated using the following equation:
      (15)        exp -
                                -  0.12 \
                                -  0.12 I
     where:      AC^  =  the relevant annualized costs associated with ERi
                 AC2  =  the relevant annualized costs associated with ER2
                     =  3.09 kg  VOC/Mg product
                     =  0.2 kg VOC/Mg product
     The annualized  costs were grouped  as  follows:  (a)  capital related
 (maintenance, taxes, insurance,  administration,  and capital  recovery
 charge, (b) labor,  (c)  pumping electricity and make-up  coolant, and (d)
 refrigeration coolant.  Table  F-7  summarizes the costs  and resulting
 exponents.
     Using the exponents  from  Table F-7 and  the  costs from Table F-6,
 the following cost-effectiveness equation was developed:
     (16)
(23,870 - 8>
CE =
                            8,170 + (15>770 .

                                     t 15 - (77,485 - 2.
                                                           - 2,085
                                                                            1.46
     where:   (23,870 - 8,170)
                       / ER - 0.2 \
                       13.09 - 0.2J
= Incremental maintenance,
  taxes, insurance, etc.
  costs associated with an
  uncontrolled emission
  rate (ER) higher than
  0.2 kg VOC/Mg product, $
                               F-15
-------
Table 7.  EXPONENTS  USED  FOR CONDENSER WITHIN
          LINE ANALYSIS,  $1,000/Mg and $2,000/Mg

Item
Maintenance
Taxes, Insurance,
Administration, and
Capital Recovery
Labor
Pumping Electricity
and Make-up Coolant
Refrigeration
Electricity
Recovery
Credit

Emission
Rate
3.0^
0.2
3.09
0.2
1
V ! , • ,„ 	 '•
3.09
0.2
3.09
0.2
3.09
0.2
i " « " T

Annuali zed
Cost Exponent
23,870
0.297 =0.3
8,170
15,770
0.742 = 0.75
1,080
2,875 6.455 =0.46
555
2,310 1.393 = 1.4
15
77,485 1.00 = 1.00
2,085
• 	 l.l iilr „,! 	 isi, ;".", • .: 	 .,11.1 	 !il, .. ,;'!<;. . .







                       F-16
                                     ,.;„«!' • .J1.
-------
                                        8,170 = Minimum maintenance, taxes,
                                                etc. costs at 0.2 kg VOC/Mg
                                                product, $
             (15,770 -



                (2,875 - 555)
1,080)/ ER - 0.2 V
      \3.09 - 0.2/
      X    1,080 '
   = Incremental  labor costs, $

   = Minimum labor costs, $
      / ER - 0.2 \0-46 = in
      13.09 - 0.2J       el
       '          '        r-r\
   = Incremental  pumping
       ectricity  and make-up
     coolant costs, $
                                          555 = Minimum pumping electricity
                                                and make-up coolant costs, $
                  (2,310 - 15)
       / ER - 0.2 \1-'
       \3.09 - 0.2J
   = Incremental  refrigeration
     costs, $
                  (77,485 - 2
      ,085)/ ER - 0.2\
           \3.09 - 0.2J
15 = Minimum refrigeration costs, $

   = Incremental  recovery credit, $
                                        2,090 = Minimum recovery credit, $
                             106
         .21/ ER - 0.2\
            G.09 - 0.2l
   = Incremental  emission
     reduction, Mg
                                         2.94 = Minimum emission
                                                reduction, Mg

                                           ER = Uncontrolled emission
                                                rate, kg/Mg

     Uncontrolled emission rates equivalent to $l,000/Mg and $2,000/Mg
were determined by trial and error, substituting different emission

rates into the above equation (16) until a cost effectiveness of $l,000/Mg

(or $2,000/Mg) was obtained.  Table F-8 summarizes the uncontrolled

emissions rates for all styrene-in-air emission calculations.
                                  F-17
-------
            Table F-8.   STYRENE-IN-AIR UNCONTROLLED  EMISSION
                           RATES EQUIVALENT TO $l,000/Mg,
                           $2,000/Mg,  and $3,000/Mg
Within Line

Across Line
                               Uncontrolled Emission Rates,  kg/Mg
                         $l,OQO/Mg
0.4454

0.1983
                 $2,000/Mg
                 $3,000/Mg
0.2903

0.1694
0.1921

0.1561
                                F-18
-------
                                    TECHNICAL REPORT DATA
                             (Please read Instructions on the reverse before completing)
 EPA-450/3-83-008
                                                            3. RECIPIENT'S ACCESSION NO.
 ». TITLE AND SUBTITLE
 Control  of Volatile Organic  Compound Emissions from
 Manufacture of High-Density  Polyethylene, Polypropylene
 and  Polystyrene Resins
              5. REPORT DATE
                  November, 1983
              6. PERFORMING ORGANIZATION CODE
                                                            8. PERFORMING ORGANIZATION REPORT NO
 I. PERFORMING ORGANIZATION NAME AND ADDRESS
 Pacific  Environmental Services,  Inc.
 1905 Chapel  Hill  Road
 Durham,  NC   27707
              10. PROGRAM ELEMENT NO.
              11. CONTRACT/GRANT NO.


                  68-02-3511
    'ONSORING AGENCY NAME AND ADDRESS
 U. S. Environmental Protection  Agency
 Office of Air Quality Planning  and  Standards
 Research Triangle Park, North Carolina 27711
              13. TYPE OF REPORT AND PERIOD COVERED

              14. SPONSORING AGENCY CODE
                NOTES
      Control  techniques guidelines  (CTG) are issued  for the control of volatile

 organic compounds (VOC) from certain  polymer manufacturing plants to inform

 Regional,  State,  and local air pollution control agencies  of reasonably available
 control technology (RACT) for development of regulations necessary to attain  the

 national ambient  air quality standards  for ozone.  This document contains information

 on VOC emissions  and the costs and  environmental impacts of RACT in polypropylene

 liquid-phase  process plants, high-density polyethylene  slurry process plants  and
 polystyrene continuous process plants.
 7.
                                KEY WORDS AND DOCUMENT ANALYSIS
                  DESCRIPTORS
   Air Pollution
   Volatile Organic Compounds
   Polymers
   Resins
   Polyethylene
   Polypropylene
   Polystyrene
                                              b.lDENTIFIERS/OPEN ENDED TERMS
 Air Pollution  Control
                                                                         c.  COSATI Field/Group
                                              19. SECURITY CLASS (ThisReport}
                                               Unclassified
  Unlimited
2O. SECURITY CLASS (Thispage)
 Unclassified
                           21. NO. OF PAGES

                                 302
                                                                         22. PRICE
EPA Form 2220-1 (Rev. 4-77)   PREVIOUS EDITION is OBSOLETE
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