United States Industrial Environmental Research EPA-600/2-80-075e
Environmental Protection Laboratory April! 980
Agency Research Triangle Park NC 27711
Research and Development
Assessment of
Atmospheric Emissions
from Petroleum Refining:
Volume 5. Appendix F
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EPA-600/2-80-075e
April 1980
Assessment of Atmospheric
Emissions from Petroleum Refining:
Volume 5. Appendix F
by
R G Wetherold
Radian Corporation
P.O Box 9948
Austin, Texas 78766
Conuact No 68-02-2147, Exhibit B
Program Element No 1AB604
EPA Project Officer: Bruce A. Tichenor
Industrial Environmental Research Laboratory
Office of Environmental Engineering and Technology
Research Triangle Park, NC 27711
Prepared for
US ENVIRONMENTAL PROTECTION AGENCY
Office of Research and Development
Washington, DC 20460
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DISCLAIMER
This report has been reviewed by the Industrial and Environmental Research
Laboratory, U.S. Environmental Protection Agency, and approved for publication.
Approval does not signify that the contents necessarily reflect the views and
policies of the U.S. Environmental Protection Agency, nor does mention of trade
names or commercial products constitute endorsement or recommendation for use.
ii
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APPENDIX F: REFINERY TECHNOLOGY CHARACTERIZATION
TABLE OF CONTENTS
Page
1.0 INTRODUCTION 1
2.0 REFINERY SIMULATION AND CLASSIFICATION 2
2.1 Single Representative Refinery Model 3
2.2 EPA/API Refinery Classification and Proposed
Models . 9
2.3 Refinery Cluster Models 39
2.3.1 Model Development 40
2.3.2 Calibration of Models 46
2.3.3 Utility of the Cluster Models 48
2.4 The Development of Refinery Plot Plans 58
2.4.1 Description of Plot Plans 59
2.4.2 Utility of the Refinery Plot Plans. ... 66
2.5 The Production of Organic Chemicals at
Refineries 67
2. 6 Summary : . . . . 73
2.7 References for Section 2 74
3.0 REFINERY RAW MATERIALS AND PRODUCTS 76
3.1 Characteristics of Raw Materials 76
3.1.1 Crude Oil 76
3.1.2 Other Raw Materials 91
3.2 Characteristics of Refinery Products 94
3.2.1 Final Products 94
3.2.2 Intermediate Products 117
3.3 References for Section 3 119
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TABLE OF CONTENTS (Continued)
4.0 REFINERY PROCESS TECHNOLOGY 121
4.1 Separation Processes 124
4.1,1 Atmospheric Distillation 125
4.1.2 Vacuum Distillation 140
4.1.3 Aromatics Extraction . 148
4.1.4 References for Section 4.1 154
4.2 Thermal Operations. 157
4.2.1 Delayed Coking 160
4.2.2 Fluid Coking 174
4.2.3 Visbreaking 183
4.2.4 References for Section 4.2 ........ 190
4.3 Cracking Operations 192
4.3.1 Catalytic Cracking 193
4.3.2 Hydrocracking 212
4.3.3 References for Section 4.3 227
4.4 Hydroprocessing 230
4.4.1 Hydrorefining 230
4.4.2 Hydrotreating 243
4.4.3 References for Section 4.4 262
4.5 Conversion Processes 263
4.5.1 Catalytic Reforming 266
4.5.2 Alkylation 288
4.5.3 Isomerization 308
4.5.4 Hydrodealkylation 320
4.5.5 References for Section 4.5 327
4.6 Gas Processing 330
4.6.1 Gas Treating/Cleaning 330
4.6.2 Product Separation/LPG Production 333
4.6.3 References for Section 4.6 334
iv
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TABLE OF CONTENTS (Continued)
4.7 Other Processes 336
4.7.1 Asphalt Processing/Production 336
4.7.2 Lube Oil Processing/Production 337
4.7.3 Blending Operations 344
4.7.4 Hydrogen Production 346
4.7.5 Sulfur Recovery 355
4.7.6 References for Section 4.7 356
4.8 Waste Treatment 358
4.8.1 Blowdown/Flare Systems 358
4.8.2 Wastewater Systems. 366
4.8.3 Sludge and Solids Treatment/Disposal. . . 395
4.8.4 References for Section 4.8 402
5.0 CONVERSION FACTORS 404
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LIST OF TABLES
Table Title Page
F2-1 Comparison of "Representative" Refinery
Product Slate with Total Actual U.S.
Production 6
F2-2 Refinery Process Unit Capacities
"Representative" Compared to Average
U.S. Refineries 7
F2-3 Subcategorization of the Petroleum Refining
Industry Reflecting Significant Differences
in Wastewater Characteristics 12
F2-4 Type "A" Refineries - Capacities as of
January 1, 1979 13
F2-5 Type "B" Refineries - Capacities as of
January 1, 1979 23
F2-6 Type "C" Refineries - Capacities as of
January 1, 1979 32
F2-7 Type "D" Refineries - Capacities as of
January 1, 1979 34
F2-8 Type "E" Refineries - Capacities as of
January 1, 1979 36
F2-9 Total Refining Capacity as of January 1, 1979 . 38
F2-10 Refineries Simulated by Cluster Models 42
F2-11 PAD Districts and Bureau of Mines Refining
Districts 43
F2-12 Summary of Major Refinery Processing Units. . . 45
F2-13 Louisiana Gulf Cluster Model 49
F2-14 Type and Number of Potential Emission Sources
for Four Refinery Cases 61
VI
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LIST OF TABLES (Continued)
Table
F2-15
F2-16
F2-17
F2-18
F3-1
F3-2
F3-3
F3-4
F3-5
F3-6
F3-7
F3-8
F3-9
F3-10
F3-11
F3-12
Title
Description and Identification Number of
Modules in a Large Existing Refinery
Product Slate for Large Capacity Existing
Refinery
Chemical Production at Refinery Associated
Sites
Organic Chemicals Produced at Refinery
Associated Sites.
Ultimate Analysis of Crude Oils, Weight
Percent
Geographic Classification of Crudes
General Properties of Crudes According to
Geographic Locations
Boiling Ranges of Typical Crude Oil
Fractions
Boiling Ranges of Typical Crude Oil
Fractions
Distribution of Hydrocarbon Types in Some
Petroleum Fractions
Potentially Hazardous Hydrocarbons in Crude
Oil
Trace Metals Found by Spectrographic Analysis
of the Ash from Crude Oil
Trace Element Contents of Some Crude Oils by
Neutron Activation Analysis
Principal Applications of Catalyst Materials. .
Major Chemicals Used in Refining and Their
Principal Uses
Compositions of Typical Refinery Gases
Page
63
65
69
70
77
79
80
83
83
85
86
90
90
92
95
101
Vll
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LIST OF TABLES (Continued)
Table Title Page
F3-13 Approximate Distribution of Hydrocarbons in
Jet Fuels 103
F3-14 Main Components of Gasoline 105
F3-15 Characteristics of Three Grades of Distillate
Heating Oils 107
F3-16 Characteristics of Two Residual-Type Heating
Oils 107
F3-17 Typical Analyses of Industrial Fuel Oils . . . 109
F3-18 Properties of Three Industrial Solvents for
Which Specifications Have Been Prepared by
ASTM Committee D-2 on Petroleum Products and
Lubricants Ill
F3-19 Comparison of Major Wax Types Produced in
the United States 114
F3-20 Major Processes and Their Major Feedstocks . . 118
F4-1 Typical Boiling Ranges for Crude Oil
Fractions 125
F4-2 Typical Emissions from Atmospheric
Distillation Unit Process Heaters 133
F4-3 Estimated Fugitive Non-Methane Hydrocarbon
Emissions from a Typical Crude Distillation
Unit 134
F4-4 Estimated Composition of Non-Methane
Hydrocarbon Fugitive Emissions from a Crude
Distillation Unit 139
F4-5 Typical Emissions from Vacuum Distillation
Unit Process Heaters 145
F4-6 Estimated Fugitive Non-Methane Hydrocarbon
Emissions from a Typical Vacuum Distillation
Unit 147
Vlll
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LIST OF TABLES (Continued)
Table Title Page
F4-7 Estimated Fugitive Non-Methane Hydrocarbon
Emissions from a Typical Aromatic Extraction
Unit 153
F4-8 Estimated Composition of Fugitive Non-Methane
Hydrocarbon Emissions from An Aromatics
Extraction Unit 155
F4-9 Typical Emissions from Delayed Coking Unit
Process Heaters 170
F4-10 Estimated Fugitive Non-Methane Hydrocarbon
Emissions from a Typical Delayed Coking Unit. . . 171
F4-11 Estimated Composition of Fugitive Non-Methane
Hydrocarbon Emissions from a Delayed Coking
Unit .173
F4-12 Process Conditions for Fluid Coking and
Flexicoking 179
F4-13 Estimated Fugitive Non-Methane Hydrocarbon
Emissions from a Typical Fluid Coking Unit. . . . 182
F4-14 Typical Visbreaking Charge and Product
Properties 185
F4-15 Visbreaker Operating and Utility Information. . . 186
F4-16 Typical Emissions from Visbreaking Unit
Process Heaters 187
F4-17 Estimated Fugitive Non-Methane Hydrocarbon
Emissions from a Typical Visbreaking Unit .... 189
F4-18 Domestic Catalytic Cracking Capacity, 1979. . . . 193
F4-19 Typical Operating Conditions for Fluid
Catalytic Cracking 200
F4-20 Domestic FCC Regeneration Techniques
(August 1978) . 201
F4-21 Energy Recovery Techniques for FCC
Regenerators 205
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LIST OF TABLES (Continued)
Table Title Page
F4-22 Emission Rates from FCC Regenerators, Before
and After CO Boiler 207
F4-23 Summary of CO Emissions from Various
Regeneration Techniques 208
F4-24 Typical Emissions from Catalytic Cracking
Unit Process Heaters 209
F4-25 Estimated Fugitive Non-Methane Hydrocarbon
Emissions from a Typical Catalytic Cracking
Unit .' . 211
F4-26 Estimated Composition of Fugitive Non-Methane
Hydrocarbon Emissions from a Fluid Catalytic
Cracking Unit 213
F4-27 U.S. Kydrocracking Capacity, 1979 214
F4-28. Licensed Hydrocracking Processes 219
F4-29 Hydrocracking - Typical Processing Conditions
and Utility Data 221
F4-30 Typical Emissions from Hydrocracking Unit
Process Heaters 224
F4-31 Estimated Fugitive Non-Methane Hydrocarbon
Emissions from a Typical Hydrocracking Unit. . . 226
F4-32 Licensors of Residual Hydrodesulfurization
Processes 232
F4-33 Residual Hydrodesulfurization 235
F4-34 Heavy Gas Oil and Middle Distillate
Hydrodesulfurization Processes 237
F4-35 Typical Emissions from Gas Oil
Hydrodesulfurization Unit Process Heaters. . . . 241
F4-36 Estimated Fugitive Non-Methane Hydrocarbon
Emissions from a Typical Gas Oil
Kydrodesulfurization Unit 242
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LIST OF TABLES (Continued)
Table Title Page
F4-37 Estimated Composition of Fugitive Non-Methane
Emissions from a Gas Oil Hydrosulfurization
Unit 244
F4-38 Light Distillate Hydrodesulfurization Process . . 246
F4-39 Processes for Olefin/Aromatics Saturation .... 250
F4-40 Hydrotreating Processes for Oils and Waxes. . . . 253
F4-41 Product Yields from Lube Oil/Wax Hydrotreating. . 254
F4-42 Typical Emissions from Hydrotreating Unit
Process Heaters 258
F4-43 Estimated Fugitive Non-Methane Hydrocarbon
Emissions from a Typical Hydrotreating Unit . . . 259
F4-44 Estimated Composition of Fugitive Non-Methane
Emissions from a Hydrotreating Unit 261
F4-45 Regeneration Schemes Used for U.S. Reforming
Capacity 279
F4-46 Operating Conditions for Catalytic Reforming. . . 283
F4-47 Typical Emissions from Catalytic Reforming
Unit Process Heaters 286
F4-48 Estimated Fugitive Non-Methane Hydrocarbon
Emissions from a Typical Catalytic Reforming
Unit 287
F4-49 Estimated Composition of Fugitive Non-Methane
Emissions from a Catalytic Reforming Unit .... 289
F4-50 U.S. Alkylation Capacity, 1979 291
F4-51 Operating Information for Sulfuric Acid
Alkylation 299
F4-52 Operating Information for Hydrofluoric Acid
Alkylation 302
xi
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LIST OF TABLES (Continued)
Table Title Page
F4-53 Typical Emissions from Alkylation Unit
Process Heaters 304
F4-54 Estimated Fugitive Non-Methane Hydrocarbon
Emissions from a Typical Sulfuric Acid
Alkylation Unit 305
F4-55 Estimated Composition of Fugitive Emissions
from an Alkylation Unit 307
F4-56 Octane Numbers of Paraffin Hydrocarbons 308
F4-57 Current U.S. Isomerization Capacity (1979) . . . 309
F4-58 Thermodynamic Equilibria of Butane, Pentane,
and Hexane Mixtures 311
F4-59 Operating Conditions for Paraffin
Isomerization Processes 316
F4-60 Typical Emissions from Isomerization Unit
Process Heaters 317
F4-61 Estimated Fugitive Non-Methane Hydrocarbon
Emissions from a Typical Butane Isomerization
Unit 319
F4-62 Typical Emissions from Hydrodealkylation
Unit Process Heaters 325
F4-63 Estimated Fugitive Non-Methane Hydrocarbon
Emissions from a Typical Hydrodealkylation
Unit 326
F4-64 Estimated Fugitive Non-Methane Hydrocarbon
Emissions from a Typical Hydrogen Production
Unit 350
F4-65 Estimated Composition of Fugitive Non-Methane
Hydrocarbon Emissions from a Hydrogen
Production Unit Utilizing Naptha as a
Feedstock 352
xii
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LIST OF TABLES (Continued)
Table Title Page
F4-66 Utility Requirements for the Production of
F4-67
F4-68
F4-69
F4-70
F4-71
F4-72
F4-73
F4-74
Emissions from Smokeless Flares
Fugitive Emission Factors for Various
Refinery Fittings
Typical Waste Stream Characteristics of
Various Process Operations. .
Classification of Refinery Wastewater
Treatment Processes
Refinery Utilization of Wastewater Treatment
Processes
Typical Efficiencies of Oil Separation Units. .
API Separation Hydrocarbon Emissions Factors. .
Sources and Characteristics of Refinery
Solid Wastes
365
375
381
381
385
396
398
Xlll
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LIST OF FIGURES
Figure Title Page
F2-1 Block Flow Diagram for a Representative
U.S. Refinery 5
F2-2 Louisiana Gulf Cluster Model Calibration .... 50
F2-3 Texas Gulf Cluster Model Calibration 51
F2-4 Small Midcontinent Cluster Model Calibration . . 52
F2-5 Large Midwest Cluster Model Calibration. .... 53
F2-6 West Coast Cluster Model Calibration 54
F2-7 East Coast Cluster Model Calibration . 55
F2-8 Arrangement of Refinery Process and Auxiliary
Modules for a Large Existing Refinery 62
F4-1 Atmospheric Distillation 127
F4-2 Vacuum Distillation 141
F4-3 Process Flow Diagram - Delayed Coking 163
F4-4 Process Flow Diagram - Fluid Coking 176
F4-5 Process Flow Diagram - Flexicoking . 178
F4-6 Process Flow Diagram - Visbreaking 184
F4-7 Fluid Catalytic Cracking Unit (FCC) 195
F4-8 Thermofor Moving-Bed Catalytic Cracker (TCC) . . 196
F4-9 FCC Unit Power Recovery System 204
F4-10 Process Flow Diagram - Single Stage
Hydrocracker ..... 217
xiv
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LIST OF FIGURES (Continued)
Figure Title Page
F4-11 Process Flow Diagram - Two Stage Hydrocracker. . 218
F4-12 Residual Hydrodesulfurization 233
F4-13 Heavy Gas Oil Hydrodesulfurization 238
F4-14 Light Distillate Hydrodesulfurization 248
F4-15 Saturation of Aromatics/Olefins 251
F4-16 Lube Oil Hydrotreating 255
F4-17 Process Flow Diagram - Catalytic Reforming . . . 267
F4-18 Process Flow Diagram - Catalytic Reforming . . . 273
F4-19 Process Flow Diagram - H2SO., Alkylation
Cascade Auto-Refrigeration System 292
F4-20 Process Flow Diagram - H2SO,, Alkylation
Stratco Effluent Refrigeration System. ....... 293
F4-21 Process Flow Diagram - HF Alkylation ...... 298
F4-22 Process Flow Diagram - Butane Isomerization. . . 313
F4-23 Process Flow Diagram - Liquid Phase
Isomerization 315
F4-24 Process Flow Diagram - Hydrodealkylation .... 322
F4-25 Hydrogen Production by Steam Reforming 348
F4-26 Hydrogen Production by Partial Oxidation .... 354
F4-27 Typical Refinery Slowdown System 360
F4-28 Example of Refinery Stream Segregation . . . . . 380
F4-29 Schematic Diagram of Typical API Separator . . . 383
F4-30 Recycle-Flow Pressurization Scheme for a
Dissolved Air Flotation Unit 387
XV
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SECTION 1
INTRODUCTION
This appendix contains a characterization of refinery
technology. The information contained here is useful as the
basis for an environmental assessment of the petroleum refining
industry.
Section 2 of this appendix contains information on
classifying and characterizing petroleum refineries. Four types
or sets of refinery models which could be used to simulate the
entire United States refining industry are included.
A description of the materials used in the refining
process is given in Section 3. Included are discussions on the
characteristics of crude oils, other raw materials, intermediate
products, and final products.
Section 4 of this appendix contains detailed informa-
tion on a variety of refining processes. This information
includes a description of the purpose of the process and a dis-
cussion of current process technology. In addition, estimates
of atmospheric emissions which result during the operation of
each process are also included.
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SECTION 2
REFINERY SIMULATION AND CLASSIFICATION
To develop an environmental assessment of refineries,
it is necessary to study refineries and refining processes in
some detail. Representative or model refineries are generally
very useful when performing this type of study. These models
can include refinery flow diagrams, material balances, process
types, operating conditions, product slates, plot plans,
emissions sources, etc.
In developing refinery models for the simulation of
the entire United States refining industry, a considerable com-
promise must be made between the number of refinery models used
in the simulation and the accuracy of the simulation. On Jan-
uary 1, 1979, there were 288 operating refineries in the U.S.1
The U.S. refining industry could be precisely and accurately sim-
ulated with 288 refinery models, each representing an existing,
operating refinery. Such a large set of models would obviously
be unmanageable for study purposes. A reduction in the number
of refinery models provides a more manageable system, but the
ability of this system to accurately simulate the refining
industry is reduced.
Four types or sets of models which have been or could
be developed are discussed in this document. In Section 2.1, a
single representative refinery with a fixed process configura-
tion is described. The distribution of products in the product
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slate of this model refinery is the same as the distribution in
the total U.S. refinery production.
Another set of five refinery models could be
developed from the five EPA refinery categories,2 which are
modifications of the five refinery classes proposed by the
American Petroleum Institute (API). This approach is discussed
in Section 2.2.
In Section 2.3, a set of "cluster" refinery models
developed by Arthur D. Little, Inc.3 is discussed. These models
were developed to simulate the U.S. refining industry in order
to study the impacts of SOX emissions control and gasoline lead
regulations on the industry. The cluster models are quite so-
phisticated and are described in some detail.
Pacific Environmental Services'* has developed
refinery plot plans for four refinery cases varying from a
small existing plant to a large new refinery. Each case has a
detailed plot plan and a census of fugitive emission sources
for every refinery unit. The cases are discussed in Section
2.4.
2.1 SINGLE REPRESENTATIVE REFINERY MODEL
It is often customary to propose a "typical" or
"representative" refinery model for purposes of illustration,
discussion, or crude refinery simulation. Flow diagrams of the
representative refinery can be used to describe the different
types of refineries, the arrangement and integration of refinery
processing units, and the origins and fates of the various
product and process streams.
-------
An example of this type of refinery model is a
representative U.S. refinery model that was defined for process
analysis and stream characterization purposes.5 The flow
diagram of this plant is presented in Figure F2-1. The product
slate for this refinery is given in Table F2-1, and the capaci-
ties of the various individual processing units within the
refinery are shown in Table F2-2. This model refinery was
developed under the following assumptions:
A. The refinery capacity was 100,000 barrels/day.
B. The process units shown on the flow diagram
are those in common use in the refining
industry.
C. The capacities of the process units relative
to the crude feedstock agree with the average
1978 capacities of all U.S. refineries, as
shown in Table F2-2.
D. The refinery product distribution is reasonably
consistent with that of the entire refining
industry in 1974.
E. The crude feedstock was a weighted composite
of crudes from the major oil fields supplying
petroleura to domestic refineries in 1974.
In theory, the entire U.S. refining industry could be
simulated with a number of these representative model refineries.
Practically, however, there are a number of serious drawbacks
to this simulation method, and only a very superficial and
-------
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Figure F2-L. Block flow diagram for a representative U.S. refinery.
Reference 5
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TABLE F2-1.
COMPARISON OF "REPRESENTATIVE" REFINERY
PRODUCT SLATE WITH TOTAL ACTUAL U.S.
PRODUCTION
Volume Percent of Total
Refinery Products
Product
Gasoline
Kerosine
Jet Fuel,
Naphtha type
Kerosine type
Distillate Fuel Oil
Asphalt
Residual Fuel Oil
Marketable Coke
LPG
Petrochemical
Feedstocks
Other (Fuels, misc.)
TOTAL
Representat ive
Refinery Production
50.3
1.2
1.5
5.0
20.4
3.4
8.2
1.4
2.4
2.8
3.4
100.0
Total U.S.
Production (1974)
48.9
1.2
1.5
4.9
20.4
3.4
8.2
1.4
2.4
2.8
4.9
100.0
Source: References 5 and 6
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TABLE F2-2. REFINERY PROCESS UNIT CAPACITIES "REPRESENTATIVE"
COMPARED TO AVERAGE OF U.S. REFINERIES
Unit
Reformer
Fluid Cat Cracker
Hydrocracker
Coking
Asphalt
Isomerization
Alkylation
Naphtha HDS
Distillate HDS
Gas Oil HDS
Res id. Oil HDS
Volume Percent
of Crude Feedstock
Representative Average of
Refinery U.S. Refineries (1978)
24.6
28.9
5.6
1.4
3.6
0.8
5.6
20.8
11.3
3.5
0.04
21.0
27.6
4.9
6.0
4.5
0.7
5.1
20.0
12.3
5.9
0.7
Source: References 1 and 5
-------
imprecise analysis of nationwide refinery operations can be
accomplished with this technique. Some of the more obvious
deficiencies of this method are listed below.
A. The types of refinery process units are
fixed. There are, however, many refineries
which utilize fewer process modules, and a
number which employ more.
B. The configuration of the process units
within the refinery is also fixed. This
greatly limits the flexibility of the
model.
C. While the representative refinery can
efficiently process the selected
composite crude oil, it is not suitable
for refining many of the individual crudes
processed by refineries across the U.S.
D. The necessary specification of the
individual processing units and the
associated range of operating conditions
further limits the utility of this model
for representing the entire industry.
E. Total fugitive emissions cannot be calcu-
lated since the model does not provide a
means of estimating the number of individ-
ual fugitive sources within each process
unit.
-------
F. The model is very general and accurate
estimates for both hydrocarbon and non-
hydrocarbon emissions from non-fugitive
sources are difficult.
In actuality, the only major variables that can be
manipulated in this type of model are the size of the refinery
(in terms of crude feedstock) and the process operating condi-
tions (within the 'operating constraints of each of the individ-
ual processes). For the purposes of this project, the single
representative refinery model is unsuitable.
2.2 EPA/API REFINERY CLASSIFICATION AND PROPOSED MODELS
One possible means of modeling the refining industry
is to categorize the nation's petroleum refineries according to
a logical set of basic characteristics, and to develop a typical
or representative model refinery for each of the categories.
The refining industry could then be simulated with this group
of models.
API has developed a classification system for
refineries. All refineries are grouped into one of the follow-
ing five basic refinery classes.
• Class A: Topping - The principal operation
is the separation of crude oil into its
major fractions. Excluded from this class
are refineries with cracking and/or coking
processes. Hydrotreating operations may be
carried out in this type of refinery.
-------
• Class B: Topping and Cracking - In addition
Co the crude separation and hydrotreating
operations of the topping refineries, cracking
and coking processes are included. Also
included are refineries with production of
conventional refinery-associated first-
generation products and intermediates such
as BTX (benzene/toluene/xylene), alkanes,
alkynes, alkenes, and other miscellaneous
products such as sulfur, hydrogen, and coke.
Refineries which contain lube oils processing
units are not included in this classification.
• Class C: Topping, Cracking, and
Petrochemical - Refineries in this classi-
fication contain topping, cracking, con-
version, coking and, in addition, first and
second-generation petrochemical operations.
No lube oils processing occurs in this class
of refineries.
• Class D: Integrated - This classification
includes refineries with all of the Class B
operations plus lube oils processing.
• Class E: Integrated and Petrochemical -
Petrochemical manufacturing units are
included in this classification of refinery,
along with all the processes of a Class D
refinery.
The EPA has developed refinery subcategories which
are reflective of the wastewater loading with respect to
10
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refinery type, processing units, and operating severity.2 There
are five subcategories that are similar to the API classifica-
tions. The characteristics of the subcategories are summarized
in Table F2-3.
The topping Subcategory A is the same as the API
Class A. Subcategory B (Cracking) is the same as the API
Class B except that the production of first-generation petro-
chemicals is included only if these products amount to less than
15 percent of the refinery throughput. The petrochemical Sub-
category C is similar to the API Class C. Petrochemical opera-
tions of refineries in this Subcategory include production of
intermediate chemicals such as cumene, phthalic anhydride,
alcohols, ketones, styrene, etc. These chemicals are generally
considered to be second-generation petrochemicals. Refineries
which include the associated manufacture of second generation
petrochemicals belong in Subcategory C.
The lube Subcategory D is the same as the API Class D.
The integrated Subcategory E is identical to the API Class E
except for the exclusion of refineries having first-generation
petrochemical operations where the petrochemical production
capacity is equivalent to less than 15 percent of refinery
throughput.
There were 288 operating refineries in the U.S. as of
January 1, 1979.1 These refineries have been classified into
the EPA subcategories, and are tabulated by Subcategory or type
in Tables F2-4 through F2-8. Included in these tables are
company, location, crude capacity, major process unit capaci-
ties, and production capacities for each of the 288 refineries.
They were classified using information from Oil and Gas Journal,1
EPA,2 and the 1976 Directory of Chemical Producers, U.S.A.7
11
-------
TABLE F2-3,
Subcategory
SUBCATEGORIZATION OF THE PETROLEUM REFINING
INDUSTRY REFLECTING SIGNIFICANT DIFFERENCES
IN WASTEWATER CHARACTERISTICS
Basic Refinery Operations Included
Topping and catalytic reforming whether or
not the facility includes any other process
in addition to topping and catalytic process.
This subcategory is not applicable to facili-
ties which include thermal processes (coking,
visbreaking, etc.) or catalytic cracking.
Topping and cracking, whether or not the
facility includes any processes in addition
to topping and cracking, unless specified in
one of the subcategories listed below.
Topping, cracking and petrochemical operations,
whether or not the facility includes any
process in addition to topping, cracking and
petrochemical operations, except lube oil
manufacturing operations.
Topping, cracking and lube oil manufacturing
processes, whether or not the facility
includes any process in addition to topping,
cracking and lube oil manufacturing processes,
except petrochemical operations.*
Topping, cracking, lube oil manufacturing
processes, and petrochemical operations,
whether or not the facility includes any
processes in addition to topping, cracking,
lube oil manufacturing processes and petro-
chemical operations.*
*The term "petrochemical operations" shall mean the production
of second generation petrochemicals (i.e., alcohols, ketones,
cumene, styrene, etc.) or first generation petrochemicals and
isomerization products (i.e., BTX, olefins, cyclohexane, etc.)
when 15% or more of refinery production is as first generation
petrochemicals and isomerization products.
Source: Reference 2
Topping
Cracking
Petrochemical
Lube
Integrated
12
-------
TABLE F2-4. TYPE "A" REFINERIES - CAPACITIES AS OF JANUARY 1, 1979
Charge Capacity, MBPSD
Production Capacity
Cat.
Cruck-
Crude Capacity „ ,. , *"*'.
y_c! i Vacuum Thcimal Ifreeh
Company - Location EPA MIL KBPCD HBPSD Dlatlll. 0|>na. Feed
Cat. Cat. Cat.
Cat. Hydro- Hydro- Hydro- Aroouitlcs/ Coke,
Reform. Crack. Rc'lln. Treat AlVylatlon laom. Lubea Aaphalta T/D
Golden Eagle Raflnlng
Co., Canon. Calif. A
Lundiiy-Thagard Oil
Co., South Cat*.
Calif. A
KjclllUn Ring-Free
Oil Co., Signal
HIM. Calif. A
Hohavk P
Corp., Inc.•
Baktrafleld,
C.1II. A
Hcvhall Refining
Co., Inc..
Htvh.ll. Cilir. A
Road Oil Salea,
Inc., BaketldeU,
Calif. A
Sabre Refining,
Inc., BakarafUld,
Calif. A
San Joaquln
lUf Inlng Co. ,
Olltlale. Calif. A
Sunland Raflnlng
Corp., Dakerafleld,
Calif.
Aaoco Ull Co.,
Savannali, Ca.
Young Refining Corp..
Douglaivllle, Ca.
Hawaiian Independent
b*flrury, Inc.,
Oalio, Hawaii
16.5
12.0
11.6
22.1
l.J
a.6
27.0
15.0
16.0
5.0
(.2.5
2.5
17.2
10.0
12.2
22.8
If).5 8.0
1.6 —
9.0 —
28. A
15.0
19.0
5.0
65.5
2.2
2.5
2.5 2.5
3.0
1.0
1.5
11.0
11.0
11.0
-------
TABLE F2-4. Continued
Category Crude Capacity
Company - Location EPA ADL HBI'CD MBPSD
Yetter Oil Co. .
Colrwr, 111. A 1.0 1.1
Princeton Rpflntng,
Inr.. PrlncrtAn. Irid. A 4.6 4.6
Gladlcux Refinery.
Inc., Fore Wnyno,
Ind. A 10.6 12.2
Lakcton Asphalt
Rt* ( Intng, Inc. ,
Laketon, Ind. A 8.1 9.J
Hill-Amor lea Refinery
Co . . Chaniue,
KansftS A 3.1 3.3
Sonernet Refinery,
Inc.. Somerset, Ky. A 3.0 5.3
Atlas Processing Co.,
Dlv. of IVimeoll,
Shrcveport. La. A 45.0 47.4
Calumet Refining Co.,
Princeton, Ln. A 2.4 2.4
Clalborne Caroline
Co. , Lllnon, Ln. \ 6.5 6.7
Evangellnc Refining
Co., Inc.. Jennings,
La. A 4.8 5.0
Cotton Valley Solvents
(Kerr McCee), Cotton
Valley, La. A 11.0 11.2
Lilec, Inc.,
St. JamCB. La. A 20.0 21.1
Placid Refining Co.,
Port Allen, I.a. A 34.2 36.0
Cbargo Capacity, MRPSD Production Capacity
Cat.
Crack- KBPSD
Vacuum Thcimal Prcnh Cat. Hydro- Hydro- Hydro- Arunutlca/ Caka.
DUtlll. Opno. Feed Rcfom. Crack. RcEln. Trent Alkylatlon I«nn. Lube* AlphiUl T/D
1.0
1.5 — — 1.5 -- — -- — —
..
7.J -- -- — - — — — J.O —
1.8 — — — — — — — — _
1.0
9.1 — — 10.0 — — 10.0 — 1.0 2.2 0.6
4.8 — — — — — „ — j.o o.J
— J.2 — — — — — —
0.6 — — — — -- —
20.0 — — — — — — — — — —
20.'0 — — i.5 — — 6.0
Continued
-------
TABLE F2-4. Continued
Charge Capacity, MBPSD Production Capacity
Cat.
CraeV- HBPSt)
Cntr-Rory Crudr Cnpoclty vacjun Thi.ti.al Ft 'l Cat llyjto Hydro Hydro Art. It tl
Company - location EPA Am. MDPCD HBPSD nlntlll. Opns. Feed Reform. Crnck. Rcfln. Trent Alttylntlon iKom.
Atnoco Oil Co. ,
B.iUliori;, MJ. A 15.0 17.0 — — -- — — — —
Chfvron U.S.A. Inc.,
Baltimore, Md. A 13.5 14.2 13.8
Cryst.il Koftnlnj Co.,
Cjrson City, Mich. A 6.2 4.0
Ljkoalil.' Re (In Ing Co.,
Hll;i7ia!C.o. Mich. A 5.6 5.9 — — — 2.0 —
Osccola Rofininfl Co.,
Wost Branch, Hlch. A 12.5 10.0 ~ — — 1.6 — — 1.8
SoutMnnd Oil Co. ,
Lurahorton, Minn. A 5.8 6.6
Southland Oil Co.,
Sanilorsvllle, MlM». A 11.0 12.5 6.9 -- — —
Southland OH Co.,
Yii^oo City, Mis*. A 3.8 4.5 2.5
Carlbon Four Corners,
Inn., Klrl.ind,
Sow Moxlco A 2.4 2.5
Southern Union
Ref Inlng Co. ,
l.ovlnglon, N.M. A 36.0 37.0
Southern Unlnn
Refining Co.,
Monument, M.M. A 5.0 3.2 — — — — —
Hoill.l.itid Oil &
RoCl'.ilng Co.,
Dickinson, K.D. A 5.0 5.3
FlnJlny, Ohio A 20.4 21.0 8.0
' Coke,
Lubes AaphalO T/D
10.7
11.0
_
—
—
2.8
5.0
1.9
..
_
4.5
Continued
-------
TABLE F2-4. Continued
Clior&o Capacity, KUPSO Production Capacity
Cat.
Crock- MBPSD
Ing, Cat. Cat, Cat.
Cntcgori Crude Capacity Vncuull Thcrci«l Preah Cat. Hydro- Hydro- Hydro- ArooatlcB/ Coke.
Corp.iny - Location CfA ACL rtRPCD MDTSD Dlor.111. Opns. PccJ Reform. Crack. Hufln. Treat Alkylatlon Ii.uu. l.ubc» AiipruillB T/D
Alllfd Mm-rUle Corp.,
Strou.l. Okln. A 7.0 7.3 7.1 — — — — -- -- — — 1.0 1.8
TunV.iw.i Refining Co.,
Arm-It, Ok l.i. A 6.0 6.J
St.inJnrJ Oil Conpnny
o« C.illf. , rortlaud,
On-son A 14.0 14.7 15.0 — — — — — — — — — 8.6
Tcnnzoll Co., WolC'n
Ili-ail Division,
Xuusovllio, Pa. A 12.0 12.5 4.D — — 3.6 ~ — 7.2 ~ — — 3.9
Kunil.ill-An.illi! Dlv.,
Wltco Clu-mlcai
Corp., Bradford, Pa. A 9.0 9.5 — — — 2.0 — ~ 2.5 ~ — 4.7
Vilvollnr Oil Co.,
Division of Aahland
Oil, Inc., Fvtcdon,
Pa. A 6.8 7.0 3.6 — — — — ~ — — ~ — 1.*
LcngvUu Rrflnlng Co.,
Division of Crystal
Oil, LongvU-v, Texas A 8.8 9.0 -- — — 5.5 — 3.0 *,0 — — — ~
Dorc!iC3LLT Rnflntng
Co., hltltc Deer,
Texas A 1,0 1.0 — — — 1.0 — — ~ —
Eddy Refining Co.,
Hou.Htuu, Tru.TS A 1.7 >. J — — — — —
Flint Chcmtcnl Co.,
Sjn Antonio, TrK.13 A 1.2 1.4 -- — — — — — — -- — -- --
Call SI.ucs Oil 6
Refilling Co., Corpus
Chrlrtl, Texas A 12.5 13.2
Howe 11 Corp..
Snn Antonjo, Tcxaa A 3.0 4.0 — — — 1.3 — — — — 0.0
Continued
-------
TABLE F2-4. Continued
Charge Capacity, HRPSD
Cat.
CcocV-
inp Cat Cat Cat
Category Crude Capacity v 1 F h C t II d o- H d 0- llvd o-
Coapany - location CPA ADI. HBPCD MDP3D DletiU. Opna. Vfci Reform. Crack. Rcfln. Treat Alky lot Ion
Sector Refining Co.,
Tucker, Texas A 9.7 10. 0
C.irhonlt Refinery. Inc.
Ucarnc. Texas A 10.0 11.0
PrlJi- Refining, Inc.
Ablli-ne, Texas A 20.5 36.5 12.0
fouth Hanpton
Hot InlnK Co. ,
Sllibcc, Tcxon A 20.5 22.5 — -- -- 4.0
Toft^ro Petroleum
Corp.. CArrlzo
Si-tint*, Tcxu« A 26.1 27.5 . — — — 4.0
Tex.ico, Inc . ,
T«>x.»s Asphalt &
Kef Inlnijt Co. ,
F.UI..-HI. Texaa A 5.0 6.0
Hunt Oil Co.,
Tusc.iloosj. Ala. A 29.0 30.0 15.0 — — 5.5 — 9.0 11.5
Vulcan Refining Co.,
Cordjna. Aln. A 10.6 11.4 — -- — 5.5 — — . ~
Vjittcr Asphalt Corp..
Tnic.llnrxi, Ala A 2.9 3.0 — — — — — -- — — i
Atlantic Rlchflrlcl Co.
North Slope, Aluaka A 13.0 13.7
North role Refining,
North Pole, Alaska A 72.6 25.0
Chevron, U.S.A. Inc.
Kcnal, Alaska A 22.0 23.2
Production Capacity
MBPSO
Aromatic o/ Cok«,
Una. Lubtl Aaphalta T/D
—
—
—
—
—
— ~- 9.0 --
—
10.0
i.O
1.5
—
—
0. 3
Continued
-------
TABLE F2-4. Continued
00
Cnteftory Crudfl C
Company - Location EPA ML MBPCD
kon.it, Alaska A j».0
Arlzontt Fm'ls Corp..
Fredonla. Arizona A 6.0
Crcvi3 Oil & Refining
Stcvona, Ark. A 2.9
01 1 Co. , Norplilct ,
Ark. A 4.4
Douglas OIL Co. ,
rar.inount, Calif. A 46.5
Douglas Oil Co. .
Santa Maria, Calif. A 9.5
EJ^lngton Oil Co..
Lorif R.-orh, Cnllf. A 29.5
EdRtngton Oxnard
Refinery, Oxriord,
Calif. A 2.5
needier Oil 4
Carson, Calif. A 25.0
GolJrn Dour Dlv..
Ultco Chemical Corp.,
Olld.iK-, C.illf. A 10.5
Suuml Refining. Inc..
Tacocii, Woili. A 8.0
Chovron U.S.A. Inc.,
Seattle, Wash. A 4.b
U.S. Oil & Ri-flnlng
Co., Tacona, Unsli. A 21.4
Charge Capacity. KBl'SD Production Capacity
Cat.
Crack- HB?SD
2E- " VacuiiD Tlicrtul Freeh Cat . Hydro- Hydro- Hydro- Aronat lea/ Cok*.
ttUPSD Dial 111. Opns. Peed He form. CracV. Kef In. Trent AlVyfntfon [son. tubes AtpMlCc 7/0
40.0 — -- — 6.0 — — 6.0 — — — —
5.4 3.5 — — — — — — — -_ — —
3.0 2.0 — — — -- — — - - l.J
4.5 3.0 — — -- — -- — — — 2.0 1.5
48.0 28.0 -- — 11.5 -- — 29.0 — — — 15.0
10,0 7.8 — — -- — — — — — — 6. B
30.0 12.6 — — — — — — ~ — -- ».0 —
2.5 - - - - - - - - - - -
26.0 — — — 4.0 — — 4.0 — . — — —
11.0 9.5 — — — — — — — — 4.0 J.2
10.0 3.5 — — — — — -- 1.9 2.6 —
4.7 5.0 - - ~ «.0 -
22.5 4.8 — — 3.0 — — J.O
Contlnu«d
-------
TABLE F2-4. Continued
Charge Capacity, KBPSD
Produrtlon Capacity
Cat.
Crack-
Category Crude Capacity B- ' "' ' '
- ° ' lapactLy V,C1|I1B Thernu,i. r,c,,t\ Cat. Hydro- Hydro- Hydro- Aioratlca/ Coke.
CoTpany - Locution EPA ADL HDPCD KBPSD DlBClll. Opno. Feed Reform. Crack. Refill. Treat Alkylatlon [con. LubM Asphalt a T/D
P.-nnioll Co., Elk
Refining Dlv..
r.illlrR Rock. U.Va. A 4.9 5.2 2.5
Qinkrr Staff nil
Re f In Ing Co i p..
Scvi'Ll, U.V«.
OuAkc-r State Oil
R,-f inlng Cor|...
St. M.xry'9, W.Vd.
Mountaineer RcflnlnR
Co., Inc., LnBflrRO,
Sn^c CrccV Refining
Co., CrowLcy, Wyo.
Co.,
rn Refining
rgu. Wyo. A
T.oulst.tn.i I.Anil !•
t'.x^lornt Ion Co..
S.ir.Tlnnd, Ala. A
Moblli' B;iy R<-flnlng
Co., Chlckfls
-------
TABLE F2-4. Continued
Charge Capacity, MBPSD
Production Capucltjr
t-0
O
Cat.
Cr.ck-
HBPSD
S«L«aorjt Crude Capacity Vacuum Thermal Ftcn'h Cat. Hydro- Hydro- Hydro-
Kentucky Oil t
Rt.>t tnlng Co. , Inc. ,
Betsy Ljyuo, Ky.
Itayou SCJtc Oil Corp.,
Houston. La. A
Cjlcaslcu Refining
Ltd. , Uke Charles,
La. A
Hill IVtroleun Co.,
Krotz Springs, La. A
Har.irhon Oil Co.,
Cdryvllle, Le . A
Mt . Ally Refinery Co..
Nt. Airy. La. A
Aromatic*/ Cok«t
Company - Location EPA ADL (ID PCD MB PSD Distill. Opus. Feed Reform. Crack. Hefln. Trnat Alkylatlon ISOB. l.ubea Aiphlltn T/D
VV.it Co.ist Oil Co..
OlUi»U. CnlU. A 19.0 20.0 — — — — — — — — — fe.O
Oru tactual Oil Co..
LVnvrr. Colo. A 10.0 11.0 2.5 -- — 6.5 — — 7.0
Cary Western Co.,
Friilta, Colo. A 13.0 14.0 10.0 — -- 2.0 — ~ 2.0 — — — O.t
Scmlilolc Asplmlt
Ki! fining, Inc.,
St. M.lrks, fin. A O.e 9.1 5.0
H. T. RUIuirde, Inc..
Cro^svlllo, 111. A 0.7 0.7
Indu^trlAl fuel &
Asplult oC Indiana,
Inc.,
Ind. A
E-2 Sorvtf Refilling
In,-. . Sliullovatcr.
9.8
4.6
0.5
5.0
5.7
10.3
4.8
1.0
5.J
6.0
2.0
10.1 10.7
200.0 205.0 100.0
13.6 14.1 — '
37.5
86.5 37.5
Continued
-------
TABLE F2-4. Continued
Category Crude
Cotip.iny - Location EPA ADL MBPCD
Slirvhi-rd Oil Inc.,
Metffientau. La. A
T i S Refining Inc.,
Ji-nnlng.i. l.n . A
Dow Chenlcal U.S.A. ,
Bay City, Mich. A
Er^on R^flnlnR Inc.,
Vlck.f burg, Hlaa. A
Kcnco Refining Inc.,
Wjlf Point. Mont. A
No«Md:« Refining Co..
Tpnouah, Nevada A
ATC Petroleum, Inc.,
Ni-vlngtoii, H.H. A
Giant Industries. Inc..
F.irnlnc.lon, N.H. A
Clhro Petroleum
Produce s , Inc. ,
Albany, N.Y. A
Quaker State Oil
Refining Corp.,
Fmlenton, P«. A
Quaker State Oil
Rcf Inlnp Corp. ,
Faruors Volley, PH. A
Adobe Refining Co.,
LA Dlnnci, Texas A
P longer Refining, Ltd.,
Mxop , Texas A
Qultflnn Kef In In* Co.,
Quit man, Texas A
10.0
10.2
14.0
10.0
4.5
4.0
12.8
8.8
28.0
3.3
6.5
5.0
4.9
6.0
Charge Capacity, MOPSD
Cat.
Crock-
Capacity Vacuum Thrroiol Prcnh Cat Hvdra llvdro Hvdro
HBPSD Dleclll. Opns. Feed Reform. Crock. Rcfln. Treat Alkylittlon
.
10.0
' 10.9
22.0
10.0
4.7 — « — — — —
4.2 -- — — —
13.0 — — — — —
9.2 — — — — —
30.0
3.5 1.7 — -- 1.3 — — 1.5
6.8 2.8 — -- 1.9 — — 2.:
5.0
5.0
5.7
Production Capacity
HBPSD
Aronotlca/ Coke,
Tfcon. Lube* Aaphalta T/D
—
_
—
_ _ —
• — — — _
—
—
--
—
1.7
2.S — —
—
—
—
Continued
-------
TABLE F2-4. Continued
Catcg
Conpany • Location EPA
R.incho RcflnlnR Co.,
Don:m, Texas A
. Saber Ref intitij Co. ,
Corpus Chrlscl, Tx. A
Corpus Chrletl, Tx. A
Slgnor Refining Co.,
Three Rivrrs. Tx. A
Thriftily, Inc.,
Gr.ih.ini, Tx. A
|SJ Tin-entry Corp.,
fO lnRlt;slde, Tx. A
Unl Refining Co..
IncloMde, Tit. A
Morrison Petroleum Co.,
Wools Ct033, Utah A
C & II RcUnnry. Inc.,
Li. ok. Uyo. A
Clnclcr Park Co. ,
Os.ige. Wyo. A
Cluiiruck, Wyo. A
ATC Pttrolcura, Inc.,
Wilmington, N.C. A
t:rlckcot\ Roflnlng Co.,
Port Nochcn, Tx. A
I'nlttd Independent Oil
Oil Co., Tacona, Unsh. A
TOTAL
ory Crude
AOL OTPCO
1.2
20.0
10.0
22.8
1.3
6.5
11.9
2.5
0.2
3.9
1.0
11.9
30.0
1.0
1762.8
Charge Capacity, HI1PSD
Cat.
Crnck-
1.2 1.2 -- - -- -
21.0 — -- — — — —
10.0
24.0 1.6 -- - fl.5 — -- 8.5
2.5 -•• ••- -- »- — « — _-
6.5
12.5 — -- — — -
2.6
0.2
1.0 — — — — — — — --
12.5
32.0
1.0 — — — — — —
1849.0 409.0 — — 168.2 — 105.4 183.4
Production Capacity
MBPSD
ArooAtlco/ Cok«,
Ifcun. Lube* Anptaita T/D
„
--
1.2
. „
--
— _
..
„
--
_
.-
1.8 3B.4 167.0
Source: Reference 1
-------
TABLE F2-5. TYPE "U" REFINERIES - CAPACITIES AS OF JANUARY 1, 1979
U)
Cnapany - Location EPA ADLa KBrCD HBFSn Distill.
Beacon Oil Co.
llanlurd. Calif. 8 — 12.3 12.4
Chanpl In Petroleum
Co., Wilmington,
Calif. B — 31.2 32.5 20.0
Kxxun Co., Bi-nlcla,
Calif. B — 99.0 105.0 54.0
Culf Oil Co.,
Calif. B -- 51.5 53.8 25.0
ia-rn County Refinery,
Inc., bakcrbf U-l cl,
Calif, B — 15.9 15.6
Hob 11 oil Co.,
Tor ranee. Calll. B WC 123.5 131.1 95.0
Pnwi-rllne Ull Co. .
Sdnli- PC SiirlngH,
Calif. B — 44.1 46.0 15.0
Shu 11 Oil Co.,
WUvlngtun, Delaware B -- 108. 0 113.0 60.0
Ti-xaco, Inc.,
WlUlngtun. Cnllf. B -- 75.0 78.9
TOBCO Corp.,
Baki-mlli-ld, Calif. b — 39.5 40.0 23. 5
Unlun Oil Co. of
Calif., Lob An^ule*,
Calif. B — 108.0 111.0 83.0
JUocv ia Ull U.S. Inc.,
Coom-rci: City, Colo. B — IV. 1 18.0 7.0
Clark Oil t Rcflnlnu
Corp.. lUrcfor. 111. B — 57.0 60.0 18.0
rg L/ p . y , Ml . r ra^ y
Cot.
Crack- irorSD
Thermal Fresh Cac. Hydro- Hydro- Hydro- Aroroutlcs/ CoVe,
Opna. Toed Reform. Crack. Kcfin. Treat AlHylatiun l&oai. Litbco Adplt.iltc T/D
0.5 — 1.7
11.5 — :- — „ — — -- — — 650
26.0 49.0 24.0 23.0 23.0 34.0 11.5 — — -- 1000
20.3 13.5 22.0 11.0 — 15. 0 3.0 -- — 4.0
6.5 3.0 -- — -- 3.0
62.0 60.0 36.0 21.7 r- 60.5 10.0 -- — — 2900
11.5 1.5 — 8.0 10.0 2.7 -- -- 1.0
41.5 35.0 24.0 — — 72.4 8.6 3.8 — — .'100
48.0 28.0 35.0 20.0 13.0 20.0 4.4 — — — 1650
7.0 12.0 15.5 14.0 — 8.4 1.8 — — — 250
20.0 45.0 49.0 21.0 ~ 85.0 8.0 -- — 10. Cl
7.0 3.0 — — — 1.8
13.0 78.0 9.2 — -- 20.0 8.0
Continued
-------
TABLE F2-5. Continued
K>
Conpany - Lneat Ion
H-irathon Oil Co. ,
Koblngon, 111.
Mobil Oil Corp..
Jollct. Ul.
Texico, Inc. ,
Lavc.-nce.IIU*. 111.
Tfcxatu, Inc. ,
LocV^ort, 111.
Union Oil Co. of
Cjlll., Lenunt, 111.
Indian* PArrn Burbau
Coo[*. Allan . , Inc . ,
Ml. Vvrnon, Ind.
Rock Inland kef tiling
Cor [*. , IndlAiiai'tl In.
Ind.
Poeti:r Ki-f Incry Co. ,
£1 Dorado, Kan.
ArluanH.il) City, Han.
CRA, Inc..
Kl.lUI|.«l.ur|i. lUn.
Ot rby Ktf? In Ing Co. ,
Wlcl.Ua, ICin.
Mobil Oil Corp.,
»«K«"l'. •<•-"••
National Coup.
Be f Intiry Atmn. ,
MePticraon, Kan.
Category Crude
EPA AOL* KBfCD
B — 195.0
B LH 180.0
B — B4.0
B — 72.0
B LH 151.0
» — 21.5
B — 43.6
8 — 21.8
B — 42.5
B — 26.6
B — 25.0
B — 50.0
B — 54.2
Capacity
MEPSD
205.0
200.0
SB. 4
75.8
158.9
22.6
44.5
22.5
47.2
27.5
27.7
52.0
57.0
Charge* Capacity, HBPSI) ?r<>Jucllun Cup^cily
Cat.
Crock- Mlil'SD
Ing, Cat . Cat . Cut . ~~~— -^~—~ ~~ -~~ -~ — ~— ~~ ~ ~~~~ ^ ^— ^ ^ —
Vacuun Thcro.il Prcrjh Crt. Hydro- Hydro- Hydro- Aiu-utico/ Coke.
Dlntlll. Opna. Fred Rtform. Crack. Hcfio. Treat Alkylac ion Unm. Lubl'B Asphalts T/D
62.0 21.8 38.0 47.3 22.0 6.0 22.0 7.6
88.0 34.0 92.0 47.0 — 75.0 74.0 21.0 — — — :030
25.3 10.0 38.2 26.7 -- — 45.6 7.3 — ~ J.O
14.7 30.0 3J.3 21.1 — — 40.0 8.9 — — — JOO
55.0 21.0 55.0 31.0 -- -- 86.2 15.5 3.5 — 2.5 1050
8.0 — 7.2 '.0
,
17.0 — 17.0 8.7 -- — 13.5 4.2 — -- 5.0
8.0 — 11. 0 4.0 — -- 4.0 2.0 -- — 2.0
13.0 — 9.6 16.3 3.0 — 16.3 2.6 — — 3.0
10.0 -- 8.5 b.3 — — 7.5 2.3 — — 2.0
9.8 3.8 10.8 5.0 • -- — 5.0 3.0 -- — -- 160
18.3 4.1 21.5 22.0 — — 11.5 3.8 — — 8.0
18.0 17.0 20.0 7.0 -- — 8.0 6.0 2.0 — — 125
Cone Lnued
-------
TABLE F2-5. Continued
i-n
Category Crude
Conpany - Location EPA ADL* >B1PCD
Aehland PvCroluun Co. ,
LouUvllle, Ky. B — 25.2
Continental Oil Co.,
Lake Cl.arlca, La. B — 87.0
Oood IIO['« Re finer lea.
Inc., Hctairlr, La. b — 86.0
i;ull oil Co. -Mlljncu
Kof . , hi! U Cl.onc,
C.Q. B l.r. vt',.'i
',',!( Oil Co.,
VtnKi, La. B — 28.7
Murphy Oil Corp. ,
Xxraui, La. B — 92.5
'JI..11 Oil Co.,
llorco, La. B LC 210.0
Tennvco Oil Co. ,
Chalniftt«. L*. B — 114.0
Ttijco, Inc. ,
Convt-nt , La. B — 140.0
Miration Oil Co.,
tH-troIt, Mich. B — 65.0
Total Pec rolbua. Inc.,
Alcu. HUh. B — 40.0
Continental (Hi Co.,
Wrtn.li.ill, Hlnn. 8 — 23.5
r.och ReClntjlg Co.,
RovL-Dont, Hlnn. B — 127.3
•loitl.vc-ul Rif Inlng Co.,
Oil, Inc.. St. Paul
Park. Minn. > — 67.0
C'P"clty Vacuu»
KBPSD Dlotlll.
26.0 13.0
90.0 11.5
95.0 60.0
202. 0 71.0
2D.1
95.4 40.0
240. U 90.0
120.0 23.0
147.4 36.8
67.0 25.0
42.0
24.0 9.0
131.9 80.0
A9.0 32.0
Charge
Cat.
CracV-
Thcrtaal Frcuh
0|>ne. Feud
10.0
15.5 25. 5
65.0
16. 0 7B.U
_-
10.5
64.9 100.0
9.0 22.0
13.3 77.8
25.5
16.0
9.5
23. 0 50.0
23.0
Capacity, MHPSD
Cnt. Hydro-
ReCora. Crack.
3.0
18.5
4.5
37. 5
18. n 11.5
23.0
'.6.0 24.0
35.0 10.0
33.3
16.0
10.0
3.6
15.0
12.0
Hydro- Hydro-
Rofln. Trj.it
1.0
19.0
4.5
3)1. II '.2.0
14.4
15.5 '.'..0
25.0 29.0
24.0
61.0
12.5 li.5
5.0 10.0
}.6
45.0 42.0
20.0 20.2
Production Cnp.icUy
WPSn
Aronuitlca/ Coke,
Alkylatlon IGOIE. Lubes Asphalts T/D
3.S
4.6 0.6
2.0
2H.4 16. i -- — i.J
--
3.0
13.5 — — 10.0 360
5.0 7.0 — — J50
13.9
3.S -- — 6.7
3.0 1.0
j.,2
8.5 — -- 35.0 1300
3.3 — — 14.0
Continued
-------
TABLE F2-5. Continued
ts>
Category Crude Capacity v
COBp£ny — LOCAtlOH ZVk ""' "«'•"« annm nl-tlll
Aourad* (lima. Corp. .
Purv/l*M, Htbs. B
Amoco 0(1 Co. ,
Suxar Creek, HUs. B
Cent I
L«uril. Hone. B
tV»nt Int-nr .il Oil Co.,
Bllll.ie«, Hone. B
F.Ki/,n Co.,
Billing*, Mont. B
PMlllua Petroleum Co.,
Gieae Falls. Hone. B
Vbptco Refining Co.,
Cu'. Hank. Hnnl. B
CXA. Inc.,
Scottbtilulf, Neb. B
Chevron U.S.A., Inc.,
PiTlh JtmLii/, N.J. B
Savajo Refining Co.,
Soulli Arc<«(j. H.H. B
Slit 11 Oil Co.,
Clnlzj, N.H. B
Mobil Ull Corp.,
BuKalo, H.I. B
AOJ^O on co..
K....d.n, !i.U. B
V^ntl.iml Oil Co.,
Ullllnton, N.D. D
Aohlnnd Pclrolcua Co.,
Cjnton, Ohio R
AUI» nurv.w
30.0
— 109.0
40.4
— 52.5
-- 45.0
6.0
5.3
5.6
— 168.0
— 22.6
18.0
-- 43.0
-- 52.0
4.7
64.0
nar-ju IVIMIIJA.
31.6
111.0 40.0
42.5 14.0
56.0 17.0
46.0 18.0
6.3 2.1
6.0
6.2 2.4
176.8 96.0
23.8 4.0
19.0 7.9
44.0 18.0
53.0
5.0
66.0 33.0
Chir&e Capacity, HUPSt) Production Capacity
Cut.
Crack- HUI'SD
Therml Freeh C/jt. llyilro- Hydro- Hydro— Arocvitlcs/ Coke,
7.0 16.2 5.4 — — 5.5 5.2 — — — 3JJ
13.5 42.0 J6.0 -- — 61.5 5.0 — — G.5 SOJ
12.0 12.0 — 14.0 15.0 3.0 2.0 — 6.0
15.0 15.8 — — 33.0 3.8 5.6 — 4.S
7.0 19.2 14.5 4.9 -- 35.5 3.4 — — 3.0 310
5.1 0.6 — — 2.0 — — — 0.9
2.2 — 2.3 ' — .— 3.3
2.4 0.8
33.0 39.0 — 60.0 59.0 3.0 — — 30.0
5.6 — -- — — 1.5 — -- 2.4
7.2 6.8 — — 6.8 K5 — — 0.7
.
19.0 11.5 — — 12.5 2.8 -- — 7.5
24.0 S. 2 — -- 10.0 2.6
1.1 — 2.0 — — 1.6
25.0 11.0 — 34.5 12.0 7.0 -- — 12.0
ContlnuoJ
-------
TABLE F2-5. Continued
ro
Caioi
Coopany - Location EPA.
Gulf Oil Co.,
CU-voa, Ohio B
Gulf Oil Co.,
Toledo, Ohio B
Standard Oil Co.
of Ol.lo,
Toluilo. Ohio B
Oklahona Refining Co.,
Cyril, Okla. B
Ktr r McCtt Corp. ,
w/nncvnod, Okla. B
ll,jd»on tcflnlng Co.,
Inc. , Cuahlng,
0« 1 a . B
Ilia' *r fining, Inc.,
Oknulgt.->, Okla. B
Sun Oil Co..
Duncan, Okla. B
TI-X.ICO, Inc. ,
•Jell TuUa, Okla. B
tor;.., Arrimrc,
0*1 a. B
Atlantic Richfield Co.,
Philadelphia, Pa. B
BP Oil Corp.,
Kjrrus Hook, Pa. B
Gulf Oil Co..
1'hlladt.lpMn, P.i. B
UnltuJ |u_f Inlnj Co.,
U.irn-n, Pa. B
gory Crude Capacity
ADI? HDPCD MB PSD
-- 42.7
SH 50.3
-- 120.0
-- 14.0
50.0
H.O
— 25.0
48.5
50.0
— 61.3
EC 185.0
- 16',.0
— 208.0
42.0
44.0
51.0
126.0
14.7
51.0
19.8
24.0
50.0
52.6
64. 5
185.0
177.0
214.0
42.6
Vacuuta
Distill.
13.0
12.5
68.0
5.0
10.0
7.0
3.2
17.0
15.3
30.0
106.0
85. 0
80.0
25.0
Charge
Cot.
Crock-
ing,
Thermal Frush
Opne. Feed
18.0
19.8
11.2 55.0
6.7
11.5
4.0 7.5
8.0
12.0 25.0
6.7 20.0
21.5
„
48.0
84.6
11.5
Capacity, MUPSt)
Cat. Cut.
Cat. Hydro- llyrtro-
Rufucu. Crnrk. KctLn.
10.0 -- 5.0
11.0 — 5.5
40.7 35.0
1.1
7.5 4.5
4.5 —
—
8.0
22.2
12.0 — 20.0
5k. 0 30.0 74.0
'.0.0 25.0 48.0
52.0 — 62.0
10.0
Production Capacity
t-IBl'SD
Hydro- ArutLat Lr.u/ Coke,
11.0 4.5 — — 2.9
11.0 5.5 — — 2.0
37.0 11.3 — — 7.0 030
1.1 1.7 — — 1.6
11.5 3.5 — — 3.5
8.5 2.0 — — — 100
1.5 — — 1.8
8.0 5.8 — — -- iOO
27.8 3.3 0.5
12.0 5,0 — — 15.0
54,0 — — — 25.0
lOi.O 8.2 O.G
52.0 15.0 5.0
16.1 1.4 — — 5.0
Continued
-------
TABLE F2-5. Continued
Cliargo Capacity, tflll'SU Production CopjrlLy
Cot.
Crack- M3«SD
Ing, Cot. Cat. Cxt.
Cjtrflof]^ Crucli^asJiZ Vocuun Thcrcul Krou'h Cot. Hydro- Hydro- Hydro- Aroiuitlcs/ Coke,
Conpany - Location EPA ADI. MB PCD KBPSD Ulstlll. Opns. Feed Reform. Crock. Kafln. Tf.it Alkyloclon Ison. Lubes A^ph.ilts T/D
Delia Refining Co.,
Mvnphli, Ti-nn. B .— 42.5 43.8 12.0 — 12.5 0.3 — — 13.5 3.6 — -- J.O
tarrtiL-Nier Refining Co.,
Mr.. Pleasant, Tex. B -- 2o.O 28.5 13.0 — 10.0 '..0 -- — 10.0 2.4 — — 8.0
American Petroflno,
Inc., Port Ailliur,
Tc-x. B -- 90.0 110.0 28.0 10.0 34.0 22.0 — 30.0 22.0 2.5 4.6
Ctaaplln Ptitrolfua Co.,
Corpu. Chrl.tl, Tox. B — 155.0 159.0 52.0 — 65.0 31.3 — 50.0 33.3 17.6 4.1
Charter IntL-rnjtlonal
Oil Co., lluuatun,
Tc«. B -- 65.0 70.0 22.0 10.0 '.0.0 13.5 — 29.5 46.3 4.5 2.9 — 5.0
OO Chtvron Oil Co.,
M PiiHo, Tex. B — 76.0 80.0 26.0 — 22.0 2b.O — 18.0 25.0 5.0 2.0 — 5.0
Cruwn Central
Pet rolr.iM Corp. .
Houiitun, Tex. B -- 100.0 103.0 38.0 9.5 520.0 22.0 — — 22.0 10.0 4.0 -- -- JOO
Dliuind ^h.^mrock Oil
Te«. ' ' B — 51.5 53.5 16.5 2.5 23.0 14.0 — — 14.0 8.7 1.4 — 2.5
La(;lorlA Oil & Ca0 Co.,
Tyln, Tex. B — 29.3 29.7 — 15.0 10.0 9.5 — — 7.0 3.0 — — -- 80
Shell Oil Company
(Mn.a. Tux. B — 32.0 35.0 10.0 — 10.5 11.0 — -- 11.0 3.0 1.0
Soutliwcnttm Refining
Co., [n<:. , (Icirpu*
ChrUll, Tex. II — 120.0 122.5 36.0 — 12.0 30.0 — 18.0 59.0 4.0 6.0
Ttxoco, Inc.,
Amsrlllo, Te.. C — 20.0 21.1 — 4.4 8.9 5.6 ~ — 5.6 1.7 — — — 100
Texaco, Inc.,
P.I POBO, Ttx. B — 17.0 17.9 — 4.4 7.8 3.9 -- — 3.9 1.7 0.5 — — 100
Cout InkH'd
-------
TABLE F2-5. Continued
Chjrga
y, MDPCD
Production Capacity
Cnc.
Ctack-
OiBp.iny - Location EPA ADL" KBFCU M0PSD DlatlH. Ov>na. Fetid Hctorci. Crack. Retln. Treat Mkylatlon Icon*. Lubcn
Tex.i* City Refinery,
Inc., TetaB City,
Tei. B
Independent Rciifling
Corp., Wlnnii:, Ts. D
Wlnaton RvflnlngCo.,
Fort Worth, In. »
Aaoco OH Co..
S.iIt tikt. <-lty.
Uc-ih B
Caribou four Corncra,
on 01 1 Cf*. ,
l.jkc City.
Hu.ky OH Cn. ,
Kuril, Salt Lake,
I'tiih
Phlll!!'* rrtrolci/K
Co., Wooda CrooM,
Utah
*tUn« 1C Rh l.f itld Co.,
dirtry foinr,
Oil Corp..
fernijalv, Wavh.
Shell Ull Co.,
Ti>K«cn. Inc.,
Anacurtca, Wa
119.6 130.0 49.r
16.0 H.4
20.0 JO.5 J.S
39.0 41.S
7.1 7.4 l.C
45.0 47.4
35.0
25.0 26.0 3.8
24.0
25.0 3.2
» — 104.0 110.0 45.0
B — 71.5 75.0 13.0
B — 91.0 94.0 33.0
B — 78.0 82.1 26.3
35.0 11.0
7. J
3.4 1.7
18.0 fi.O
2.0
18.0 5.5
4.4 i.O
5.2
8.4 4.7
39.0
25.5 23.0
36.0 20.0
33.3 22.2
11.0
8.3
6.0
3.5
12.6
39.0 12.0 27.0
8.5 48.0
'.6.7
1.0
5.9
12.1
O.H
2.9
Coke.
T/D
-------
TABLE F2-5. Continued
Catei
Company - Location EPA
Hurphy Oil Corp. ,
Superior. Ulu. B
Husky Oil Co.,
Cheyenne, Uyo B
llubky Oil Co.,
Cody, Uyo. »
Little Amtrlo
Kef In Inn Co. ,
Sinclair Oil Corp. ,
f^O Sinclair, Uyo. B
0
Xuvc :a»tlc, Wyo. B
Te .',<.«. Inc.,
Canpvr, Uyo. fi
t'niivrnn U.S.A. Inc..
Bakcnflild. Calif. B
Clicvrun U.S.A. Inc . .
pacific Refining Co. ,
Ikrc.iUu, Calif. B
Ulri-back Oil Co.,
Inc.. Myooulli, 111. B
KtuT&y Cooperative Inc.,
taut Chicago, Ind. B
!lavajo Xeflnlnf) Co.,
torth Arteala, N.H. R
Platciu Inc. .
fclooafluld, N.H. 1
gory Crude
AOL* KBPCD
40.0
24.2
— 10.8
-- 24 . 5
— 49.0
— 10.5
— 21.0
26.0
85. 0
1.8
— 126.0
6.0
-- 12.9
Capac Ity „
KB PSD DUtlLl.
46.8 20.5
25.2 14.0
11.3 t.j
25.8 8.6
50.0 16.1
11.0
22.1 10.5
27.4
46.3 25.0
89.5
1.8
140.0 70.0
6.3
14.0
Cliurgo Capacity. KUPSD I'ruductlon Cupnclty
Cut.
Crack- MarSD
Thurtoal Fresh Cit. Hydro- Hydro- Hyilro- Aruiuiticfi/ Coke,
Opns. Fec<) Reform. Ciar.k. R(?fln. Trc.it Alkylatlon Isoo. Lubcq Asphalts T/D
9.7 10.0 -- 5.8 10.0 1.7 — — 13.5
10.0 f.2 — — 11.1 2.8 1.5 — 3.0
3.3 1.5 -- — 3.3 0.8 — — 4.0
6.5 3.8 — — 8.8 -- -- — 2.0
17.7 9.7 — 12.2 25.0 2.2 — — 2.6
4.0 — — — — 0.9
4.4 7.8 4.4 — 4.4 4,4 — — — 1.7 125
9.8 — 6.0 — ~ 6.0
22.0 — — — 3.5 4.5 1.5 — 1.1
15.0 3.0 — 14.0
1.8
48.0 ?n.O — — 45.0 6.0 — ~ 10-4
1.5 — 7.7 — 14.4 — — —
5.0 2.3 — — 2.3
Cent inu.'J
-------
TABLE F2-5. Continued
Category Crude Capacity ..
Company - Location F.PA Ant" MBPCD HUPSD Dlr.tlll.
Karat hon Oil Co. ,
Tcxaa City, Ti. 1 — 66.0 68.0 28.0
Sun Co. . Inc .
Coipua ChrlHtl, T*. B — 57.0 60.0 10.0
TOTA1 7329.1 2665.8
*AliL AM.revUt lonn:
U; - LouUI.inn Cult Climtor
LH • Lar>{<: Hldwc.turn CLuxtur
SH • Stoall Hldcont ln«nt ClueLer
Ut: - Ucit Coast Cluiitcr
CliarRC Capacity. HUl'SO
Cot.
Crack-
Ing. Cac. Cat. Cot.
Thermal Pteoh Oat. Hydro- Hydro- Hydro-
Opno. Fcrd Hrfi>rm. Crnck. Rc-fln. Tc^ot
3B.O 8.0
7.7 25.0 24.0 — 12.5 3.2
715.0 2368.6 1637.1 334,7 829.8 2251.0
I'roduct Ion Capacity
HHI'SD
Arureatics/ Coke,
Mkylatlon iRom. Luben Asphalts T/D
11.0 2.S -- — —
7.3
404.6 37.4 — 311.1 MUO
-------
TABLE F2-6. TYPE "C" REFINERIES - CAPACITIES AS OF JANUARY 1, 1979
CO
ho
Charge Capacity, H11PSD
Cot.
Crnck-
Catcqory Crude
Conj'any - Location iiPA Adl.a KnPCU
Chrvron U.S.A. Inc.,
Kl Segimdo. Calif. C WC 405.0
Cutty Oil Co. , Inc.,
Delator. City, Del. C — 140.0
Aiuco Oil Co. .
Wood River, 111. C — 105.0
Clark Oil I Refining
Corp. , Blu« lelanu,
111. C — 66.5
C«tty Refining Co.,
tl Dorado, Kin. C SH 80.6
Chc-/ron U.S.A. Inc. ,
raflcagoula, HI... C — 280.0
Zxxon Co. ,
lindjii, K..I. C EC 290.0
Texaco, Inc.,
Vcttvllle, H.J. C — 08.0
A*til.ind Velroloua Co.,
Tunow,n.la, 11. V. C — 64.0
Sujl Co. ,
ToU'dj, Ohio C -- 125.0
Ar^oco Oil Co. .
Ttxxn City, Tx. C -- '.15.0
ivt ro-tl'cralt..l Co.,
Cu>:>ui ChrLitl. Tx. C — 135.0
Af.Tlc.m Pvtroflnj Inc.,
tig Soring, Ti. C -- 60.0
Karat turn oil Co.,
Tox-j- City, Tx. C — 66.0
Coptic Icy
UK PSD
47 r,
150
107
70
82
294
307
92
(.6
no
432
104
65
68
.3
.0
.0
.0
.0
.7
.0
.6
.0
.0
.0
.7
.0
.0
Vacuum
Dlotlll.
179
90
40
27
27
148
155
31
25
22
1V1
45
25
28
.0
.7
.0
.0
.0
.0
.0
.1
.0
.0
.0
.0
.0
.0
Tl.t-rmjl Prush Tar. HydLu- Hydro- Hydro-
0|>no. Feed Rr.Conn. Crock. Kctln. Tr^:it
54.0 52
44.0 62
38
26
11.5 31
56
135
14.4 44
23
50
33.5 184
12.0 19
10.0 24
36
.0 60.0 49.0 98.0 74.0
.0 42.0 20.0 — 110. 0
.0 12.3 — — 35.6
.0 30.5 11.0 — 20.5
.0 11.5 — 40.0 27.3
.0 CO.O 68.0 56.0 48.0
.0 19.0 — 50.0 158.0
.4 14.4 — — 44.4
.0 11.5 — 20.0 27.0
.0 41.0 ^O.U — 27.5
.0 1.14.0 42.0 — 189.0
.0 :l'i.O — 25.0 50.0
.0 20.0 — . 8.0 25.0
.0 8.0
Production Capacity
MBl'SI)
Aronklt Ice/
Alkyl.lt Ion Tsocl.
5.9 1.5
8.0 0.5
5.5
6.0
10.0 1.4
9.2 6.0
8.5
3.3
1.7 J.O
7.0 8.9
31.0 45.0
2.5 17.5
6.0 ' 7. S
11.0 :.5
' Coke.
I.ubos A:ii>l>.il 1 4 I/O
8.3 2700
1100
10.8
4.5
610
--
43.0
—
10.5
3.6
5.3 1500
0.5 500
3.0
-_ __ ._
-------
TABLE F2-6. Continued
Charge Cu;>ar.lcy, MU'.'SD Production Cai'ucltv
C.i l.
Crack-
Ing, Cat. Cat. Cat. ——
Cntt£orjr Crude Capacity Vacuum Thermal Fresh Cat. Hydro- Hydro- Hydro- AroMtlcs/ CoWo.
Coopiny - location LPA ADL1 MBPCD MBPSD Dlotlll. Opno. Peed Reform. Cr.irV. HvCLn. Tco.it AUylntlon Uoaj. l.nbcs Adpti^Ufi T/0
Phllllpn P«strolcuii Co.,
iorgit, T». C — 97.0 100.0 — — 52.0 28.5 — — 66.5 16.8 3b.»
Phllllpu Pecrolcuo Co.,
£wter.y, Tx. C — 97.0 100.0 18.0 8.0 35.5 36.0 — — 54.5 10.1 27.2
Sun Co. Irtc.,
Corpun Clirl.tl, Tx. C — 57.0 60.0 10.0 7.7 20.0 24.0 -- 12.5 3.2 7.3 — — — 235
\aoco Oil Co.,
Yorktoun, Va. C — 51.0 55.0 29.0 15.0 28.0 9.5 — — 26.5
Atlantic Rlc'.fUld
Co., Carson, Calif. C \K 180.0 186.0 76.0 84.5 56.0 33.0 19.0 — 75.0 7.2 2.i — —• 1300
(jj Howt-11 Corp.,
CO Corpus ChrUtl, T«. C — 15.0 li.8 'i.O — — 9.5 5.0 — 1U.U — 6.S
TOTAL 2BM.1 3002.1 1170.8 29'..6 973.9 084. / 240.0 309.5 1072.0 157.4 166.4 — ^65.9 SSiS
Source: Refuruncu 1
"*UI, AttlKvluC 101)01
RC - Kaxt Cutfl Clust«r
SH - Suull Hloconl tnunc CliiBLor
UC - Uo»t Coait Clu.cor
-------
TABLE ¥2-7. TYPE "D" REFINERIES - CAPACITIES AS OF JANUARY 1, 1979
Cm.
Crock-
i ....
~
opa-iy - Loco lua ZP^
To«co C:u, Tx. 0
V*ll OIL Co.,
Deer Park. Tx. 0
Icxacu, Inc. .
Port Arthur, Tx. D
Li*?LX ('r'
ABL3 iffl
j.1p
Cnpnr Iry
fU/L rlfri Uf rut 1 JL> UABkAAA.,
— «7
— 13/
— Ill
— 2t)3.
— 380
49.
LC 291
— 168.
EM 53.
— 132
— 88,
TC J25,
-- 285,
— &OA.
.0
.0
.0
.0
.0
.9
.0
.0
.8
.0
.5
,0
.0
,0
48.3
144.2
117.0
295.0
405. 0
50.0
306.3
177.0
56.0
136.0
90.0
325.0
310.0
427.4
17
81
38
95
180
17
83
51
18
32
31
103
125
l',9
.0
.0
.5
.5
.0
.0
.0
.0
.0
.0
.5
.0
.5
. 5
»>.|,
15.
17. 0 47.
42.5
94.
25.0 140.
13.0 * 16.
28.0 12>.
16.2 3V.
5.0 19.
17.0 «.
8.2 30.
27.0 114.
65.0 70.
20.0 150.
0
0
0
0
0
0
7
0
0
0
0
0
0
5
30
26
90
76
8
46
47
15
31
23
102
68
66
.0
.0
.0
.0
.0
.6
.0
.0
.0
.0
.0
.0
.0
. /
' ~~
u^-tiyn Capacity
' '
Hydvo— llyilro— llynro— Arou.it lci»
11.9 4.5
20.0 — 34.5 10.5 l.S
JO.O — 44.0
33.5 27.0 158.5 22.0 2.9
146.2 19.0 15.7
3.0 19.8 6.0
.36.0 60.0 33.0 2.3
20.0 — 59.0
20.4 4.5 6.0
. 38.0 9.7 4.7
35.0 b,6 6.6
29.0 — 206.0 12.0 20.0
50.0 198.0 7.9 15. 1:
16.7 — 155.6 16.X
n
/ Coki:,
Luboj Asphjlts T/5)
0.8 i.8
0.1 — 13JO
3.6 4.J 1B50
5.6 29.4
6.2 40.0 1200
2.S — 450
7.0 -- 1000
2.1 — 020
1.1 1.8 .200
2. 0 3.0 600
•i.l) 4.8 jOJ
8..1 — • K'OO
7.9 4.2
22.2
Continued
-------
TABLE F2-7. Continued
Chirgi! Copni;lty. MHI'SD Prodnrliiv., Ti
Cat.
Crock MM'SD
ing. Cat. Cat. Cat.
, . . . .
S!iiSJi"*J. ^L^l'^lf^^Li Vacuun ThL.tn.ol rrc»h COL. llyjro- Hydro- Hydro- Ai-uiu.it Ic»/ Ccke,
Coapany -Location EPA AOL." Mlim) MBPSD DldtlK. Opna. Pocd lljform. Crnrk. Ki'fln. Tconl Alkyl<.tl'.n Itura. I.ubcs Asph.illc T/D
CO Oil Co. ,
.pc!. Uyo. D — 47.0 48.0 15.5 — 13.0 5.8 — — 7.1 — 1.3 1.9 1.6
Oil Corp.,
fatiUboro, tl.J. D — 98.0 100.5 62.6 23.7 25.0 23.5 — — 62.6 2.J — 6.4 — 97i
A^Klord i'vtrolcum Cu.,
C.iU.-ttaburc. Ky. D — 115.8 140.0 72.7 4.0 55.0 26.5 — '.0.0 77.5 5.5 18.5 5.0 JO.O
TOfAI. 3038.0 1185.7 1171.3 351.6 995.7 6')0.9 l'.9.2 156.0 133'..! 162.2 95.6 92.2 11!.9
Soiree: Reference I
O1 ADI. Abbrevlnt Inno:
LC - LoulHtunn Cull Cluster
LK - l.arp.c mii/onc Cluster
SK • Sn.nl 1 Kldcontlnonr Clui>co
TC *• Texan Gulf Coaat CluHter
-------
TABLE F2-8. TYPE "E" REFINERIES - CAPACITIES AS OF JANUARY 1, 1979
(jO
Cltnrco Optclty, M11FSD
Cur.
CmcV,-
^i-y.a cr."^_l'.»r_"?_'t.x Vncmin, Ti,c,,nii n-TL'h c.u. Hydro- Hydro- H^'O-
ros|..in« - I.oc.illon F.p\ APl.n Miirci) Mtl'SU Distill. Ofn*. Fco.l Ki'Iorm. rrairk. Rrlln. TrcMl
SI:.- 1 1 :'U.
Mvitln. i, Cnllf. ! — IR'i.O 107.0 55.1 — 'i6.0 2'..0 20.0 JO.O 60.3
•;• V.H ...rj (cil .),
.,i.;!:::.',u', L'.i'.iC. H — 365.0 33'.. 0 220.0 -- 55.0 84.0 7i.O 125.0 85.7.
K..-..I.I C'cy. Cm. B — 76.0 H5.0 23.5 — 33.5 21.0 — — 62.0
:'.u>ui Hool:. !M. E r.C ir-5.0 IBO.O 48.0 -- 75.0 '.'..3 -- — 6if.O
i!.'.0.0 r>f>8.0 ISO.'J — U5.0 L'.S.O !1.0 19(1.0 416.5
c-,:t en.
~>..-i Aitliut. Tx. E TC 33-'.. 5 3'.?.0 157.5 30.0 120.0 65.0 15.0 «. 0 65.0
V:i:,-:i 01 !.
:icJ.-rlnn.l, TV. F. -- 120.0 12'.. 3 61.0 — 33.0 30.0 — — '.9.0
r.MAL 7.I.K.5 ?iU3.3 17.30.9 IIO.O 7'.7.5 601.3 150.0 607.0 1011.2
'.V)l. A'.^.-.t-vtnt Ions :
Production C.iparlry
mr-;i,
Arotnc lo «/ Cofc*?,
AlVyl.-itlon lion. l.ubcn Aiph-iln I/D
8.0 -- /,.5 10.5
9.2 6.0 4.3 Sl.O
».5 -- 2.9 <,.!
17.11 • -- 17.0
9.0 11.2 6.« — 1300
26.0 — 3J.8 12.0
20.0 12.'. 13.2 -- 1 ~.V>
4.2 3.9 1.« ).<,
133.0 31.5 103.5 '2.2 'j'jW
-------
Some companies reported capacities only on a stream-
day basis or only on a calendar-day basis. These capacities
were converted to the required bases using a stream factor of
0.95 for crude and vacuum units, and a stream factor of 0.90 for
all other processes.
A summary of total refining and process-unit capac-
ities for each refinery subcategory or type is given in Table
F2-9. Type A refineries are the most numerous. Howerver, more
crude is refined in Type B refineries than in any other. There
are over 130 Type A refineries, but only about 10 percent of the
crude is processed in these plants.
A representative model of each of the five refinery
types could be developed as an average of all the refineries in
each classification. The processing unit charge capacities and
production capacities of each model would correspond to the
average capacities (expressed as percent of crude) of each of
the various refinery types as shown in Table F2-9. It would be
possible to simulate the process unit capacities and product
output of the U.S. refining industry with selected groups of
these model types. An extensive amount of work would be
required, however, to calibrate these models such that they
would best simulate the operating refineries within each
refinery category.
For example, it would be difficult to arbitrarily
choose specific types of processing units and associated
operating conditions that would accurately represent actual
operations in the different types of refineries. It would be
necessary to obtain from the refining companies a considerable
amount of data concerning the operation of the individual
processing units within the specific refineries. Companies are
37
-------
TABLE F2-9. TOTAL REFINING CAPACITY AS OF JANUARY 1, 1979
CO
00
Refinery
Type
A
B
C
D
E
VOTAL
Source:
Number
of
Plants
131
111
20
17
9
288
Crude
Capacity
MflPCD
1762.8
6S96.5
2861.1
3036.0
2669.5
17235.9
KRPSD
1849.0
7247.1
3002.1
3185.7
2813.3
16097.2
Vacuum
Dlutll-
Utlon
409.1
2665.8
1170.8
1173.3
1280.9
6i>99.9
Thermal
(Jpera-
t lonn
715.0
294.6
351.6
110.0
1471.2
Charge
Cat.
Crack-
Ing.
Frcflh
Fred
2368.6
973.9
995.7
V42.5
5080.7
Capacity.
Cot.
Rofortn.
168.2
1637.1
684.7
690.9
601.3
3782.2
MBI'SU
Cat.
Hydro-
Crack.
334.7
240.0
149.2
156.0
879.9
Production Capacity
Cnt.
llydro-
Refln.
1C5.4
329.8
309.5
156.0
607.0
2C07.7
Cat.
Hydro-
Trent
183.
2251.
1072.
1334.
1011.
5851.
4
0
0
1
2
7
KHPSD
Alkylation
484.6
157.4
162.2
133.0
937.2
Aroxat Ics/
Isora. Lub.v; Asphalt
1.3
8/.4
166.4
95.6
31.5
3S2 . 7
38.4 167.
313.
265.
•;?.: 112,
10J.5 72,
234.1 929.
0
,3
.9
,9
. 3
Coke
(cony/day)
21130
8845
9595
5590
45160
Reference 1
-------
normally very reluctant to divulge this type of proprietary
information.
Linear programming models of refineries could be used
with selected crude charges, average product slates, and
specific processing units and configurations in order to develop
sizes and operating conditions of the individual processes.
Again, however, it would be difficult to ascertain the validity
of these operating conditions without an extensive amount of
data from operating refineries.
In summary, the use of five refinery models instead
of one would provide more flexibility and inherently greater
accuracy in simulating the refining industry. An extensive data
base of refinery operating data would be needed to evaluate the
validity of the models and for calibration purposes.
This model still falls short of providing the required
data for calculating fugitive emissions since the number of in-
dividual sources within each of the process units in the five
model refineries cannot be accurately estimated. However, this
model should provide a better estimate of emissions from non-
fugitive sources than the single representative refinery model;
particularly those emissions associated with specific process
units.
2.3 REFINERY CLUSTER MODELS
In some work done for the EPA, Arthur D. Little, Inc.
has defined the impact of several existing or proposed environ-
mental regulations on the U.S. refining industry.1'8'9 The
development of a method for providing a detailed simulation of
the U.S. refining industry constituted an important part of
39
-------
these studies. The acquisition and collation of an extensive
refinery data base, which was used in calibrating the refinery
models, was accomplished.
2.3.1 Model Development
In the method developed by Arthur D. Little, Inc.,3
the refining industry, as it existed in 1973, was simulated by
six individual refinery models. Each of the models, called
"cluster" models, lies in a different geographical area of the
U.S. and consists of a group of three existing operating
refineries. In terms of crude oil type, process configurations,
operating conditions, and product distributions, each group
consists of refineries that are typical of the refining indus-
try. Crude capacities of the model refineries ranged from
48,000 to 350,000 barrels/day, and each cluster model repre-
sented the average operation of the three refineries selected
to comprise each group. Representatives of the EPA and members
of a task force from the American Petroleum Institute (API) and
the National Petroleum Refineries Association (KPRA) assisted
Arthur D. Little, Inc., in the selection of the six geographical
areas and the three refineries within each area.
The two most important criteria observed in the
selection of representative regions and refineries were:
(1) each selected cluster mode was to represent, as nearly as
possible a typical and realistic refinery operation with respect
to type, size of processing units, and operating flexibility;
and (2) the crude slate, processing configuration, and product
slates for each model were to be representative of the varia-
tions peculiar to each geographical region.
-------
The six cluster models and the refineries comprising
each cluster are listed in Table F2-10. The capacities of the
refineries are shown for 1973, the calibration year, and as of
January 1, 1979. The cluster model refineries are also identi-
fied in the refinery tabulations in Tables F2-4 through F2-8.
The Bureau of Mines has grouped refining operations
into geographical refining districts which correspond to
districts designated as PAD (Petroleum Administration for
Defense) districts. These districts are briefly described in
Table F2-11. The number of cluster models selected to charac-
terize each PAD district reflects the refining capacity and the
variations in the type of available crude found in each of the
districts. For example, sufficient refining capacity and
enough crude with common characteristics are available in PAD
District I to permit characterization of that district with one
cluster model, the East Coast cluster. On the other hand, about
40 percent of the total U.S. refining capacity is contained in
PAD District III. Two cluster models were used to simulate the
refineries in this district because of its importance, and be-
cause two types of refinery configurations and crude slates were
identified.
PAD District II, containing about 28 percent of domes-
tic refining capacity, was also simulated with two refinery
cluster models, since two distinctly characteristic types of
refineries could be identified within the region. One was a
large Midwest-type refinery processing over 100,000 barrels/day
of high sulfur crudes. The other type is typified by the small
Midcontinent refinery in which 50,000 - 100,000 barrels/day of
low sulfur crudes are refined.
41
-------
TABLE F2-10. REFINERIES SIMULATED BY CLUSTER MODELS
Ni
PAD Cluster
District Identification Refineries Simulated
I East Coast Arco - Philadelphia, Pa.
Sun Oil - Marcufl Hook, Pa.
Exxon - Linden, New Jersey
II Large Midwest Mobil - Jollet. Illinois
Union - Lemon t, Illinois
Arco - East Chicago, Indiana
II Small Mldcontlnent Skelly - El Dorado, Kansas
Gulf Oil - Toledo, Ohio
Champlln - Enid, Oklahoma
III Texas Gulf Exxon - Bay town , Texas
Gulf Oil - Port Arthur, Texas
Mobil - Beaumont, Texas
III Louisiana Gulf Gulf Oil - Alliance, L,i .
Shell Oil - Norco, La.
Cities Service - Lake Charles, La.
V West Coast Mobil - Tor ranee, California
Arco-Caraon, California
Socal - El Segundo, California
1973
Crude Capacity
MB/CD
160.0
163.0
255.0
160.0
140.0
135.0
67.0
48.8
48.0
350.0
312.1
335.0
174.0
240.0
240.0
123.5
165.0
220.0
1978
Crude Capacity
MB/CD
185.0
165.0
290.0
180.0
151.0
126.0
80.6
50.3
53.8
640.0
334.5
325.0
195.9
230.0
291.0
123.5
180.0
405.0
Source: References 3 and 1
-------
TABLE F2-11. PAD DISTRICTS AND BUREAU OF MINES
REFINING DISTRICTS
PAD District
I East Coast + Appalachian No. 1
II Appalachian No. 2 + Indiana-Illinois-
Kentucky -I- Oklahoma-Kansas-Missouri +
Minnesota-Wisconsin-North Dakota-South
Dakota
III Texas Inland + Texas Gulf Coast +
Louisiana Gulf Coast + Northern
Louisiana-Arkansas + New Mexico
IV Rocky Mountains
V West Coast
43
-------
PAD District V was characterized by a single refinery
cluster model, the West Coast cluster. Since less than 5 percent
of the country's refining capacity is located within PAD District
IV (Rocky Mountains), it was not represented by a cluster model.
Four of the five EPA refinery types or categories
(which are very similar to the five API refinery classes) are
represented in the 18 refineries which make up the cluster
models. The Type A (topping) refineries are not represented.
Although there are 131 topping refineries in the U.S., they
process only about 10 percent of the total crude.
A summary of five types of major refinery processing
units and capacities in each cluster model is presented in
Table F2-12. These are compared with the average processing
unit capacities (1973/1974) in refineries located within the
same PAD Districts as the cluster models. In general, the
capacities of the cluster model(s) characterizing each PAD
district agree with or bracket the average capacities of the
respective PAD districts.
There are a few cases with an above average deviation
between the model values and the average district capacities.
In PAD District I, the East Coast cluster contains no coking
processes, even though coking capacity exists within the
district. Coking operations were purposely omitted in
order to have one cluster model without a coking unit, since
many refineries do not carry out coking operations.
The capacities of the cluster model catalytic cracking
units in PAD District II are higher than the average district
capacities for this process. However, the cluster model refin-
eries contain no hydrocracking units, so the composite of
cracking operations compares well with the district average.
44
-------
TABLE F2-12. SUMMARY OF MAJOR REFINERY PROCESSING UNITS
(Percentage of Crude Capacity)
Processing Unit
Catalytic reforming
Catalytic cracking
Hydrocracking
.p-
Oi Alkylation
Dp-layed coking
Texan
Gulf
Cluster
23
33
6
6
6
La.
Gulf PAD III
Cluster Average
IB 24
41 31
4 5
11 6
9 8
Large
Midwest
Cluoter
21
35
0
9
10
Small
Mid-
Continent
Cluster
27
39
0
9
8
Bast
PAD II Coast
Average Cluster
22 23
34 34
4 5
7 4
9 0
PAD I
Average
21
33
3
4
7
Went
Coast
Cluster
24
29
16
4
24
PAD V
Average
23
24
15
5
20
Source: Reference 3
-------
2.3.2 ; Calibration of Models
After developing the six cluster refinery models to
represent the entire refining industry, the validity of these
models was verified in reflecting regional and national refinery
operations.
2.3.2.1 Input/Output Data —
The crude oil slates for the cluster models were
developed from Bureau of Mines data and representative crudes
selected by Arthur D. Little, Inc. (ADL). The product outputs
for 1973 were also supplied by the Bureau of Mines for each of
the refineries in the various cluster models. This product
slate was then averaged over the three refineries in each
cluster in order to get a representative product slate for each
of the cluster models.
2.3.2.2 Processing Configurations—
The annual refining surveys published in Oil and Gas
Journal1°''' provided the basic data used to select the cluster
model processing configurations. Processing unit capacities as
of January 1, 1973 and January 1, 1974 were averaged to obtain
values for the calendar year 1973.
It was necessary to obtain processing conditions for
the various model configurations. Much of these data are
proprietary. To avoid publicly divulging this proprietary data,
the individual refinery processing conditions were transmitted
by the refineries to the EPA. The EPA then averaged the operat-
ing conditions for the three refineries in each cluster model,
and reported the average operating conditions to ADL.
46
-------
2.3.2.3 Other Calibration Data--
Gasoline production rates and properties such as
sulfur content, octane ratings, vapor pressure, and maximum
lead addition were obtained and compiled by the EPA. Properties
of other products such as kerosene, jet fuel, residual fuel oil,
and distillate fuel oil were also obtained and used in the cali-
bration effort.
The energy consumption for each of the refineries in
the cluster models was obtained from the Bureau of Mines.
2.3.2.4 Calibration Results--
Using linear programming models to represent the
various refinery processing units and the process integrations
within each refinery, the cluster models were simulated and
calibrated. The types of crude, inputs of materials other than
crude, processing conditions, and product slates were specified.
The amount of crude feedstock was allowed to vary in order to
close the overall material balance around the refinery. The
results were compared to Bureau of Mines and industry data.
Comparisons were made in four main areas. These areas were:
• overall refinery material balance;
• refinery energy consumption;
• processing configuration, throughputs,
and operating severities; and
• key product properties.
47
-------
As an example, the calibration results for the
Louisiana Gulf cluster model are compared to the Bureau of
Mines and industry data in Table F2-13. It can be seen that
in the areas of material balances, energy consumption, and
processing, the agreement between the model, Bureau of Mines
data and industry data is good.
Refinery process flow diagrams were developed for
each of the six cluster model refineries. They are shown in
Figures 72-2 through F2-7.
2.3.2.5 Scale-Up of Cluster Models--
After each cluster model was calibrated with Bureau
of Mines and industry data, the models were scaled-up to
represent the entire U.S. refining industry. The products and
inputs for each PAD district were determined by scaling-up the
products and inputs for the particular model or models of that
district. The scale-up was accomplished by making the gasoline
production in the cluster model(s) equal to the gasoline
production in the corresponding PAD district. The maximum
deviation of the model predictions from the total raw material
input for any single district is 6.8 percent for PAD District V.
The deviation from the total U.S. refinery raw material intake
was only 1 percent. The calibrated cluster models provide a
good simulation of the total U.S. refining industry.
2.3.3 Utility of the Cluster Models
The cluster model method for the simulation of the
U.S. refining industry appears to have a number of distinct
advantages over the two potential simulation models discussed
in Sections 2.1 and 2.2. Some of these advantages are:
48
-------
VD
TABLE F2-13. LOUISIANA GULF CLUSTER MODEL
Calibration Results (MB/CD)^
Oil and Gas
Capacity MB/SD
BOM
Data
Industry
Data
Model
Run
Material balances
Total crude intake 219.8 216.8 222.2
Energy consumption
Purchased natural gas
Total fuel consumption
Electricity MKWH/CD
Processing summary
Catalytic reforming
Catalytic cracking
Alkylation
Hydrocracking
Coking
(FOE)
(FOE)b
intake
severity RON
intake
conversion % vol.
production
intake
intake
40.5
95.0
24.0
9.5
20.0
5.4
17.2
606
32.9
92.3
90.0
70.0
17.7
6.3
19.1
5.4
17:0
744
28.3
90.0
82.2
69.6
17.5
6.6
15.8
unless otherwise noted.
Excludes catalyst coke.
Source: Reference 3
-------
'. Ml ft 10 TBi.K>rj;jl»
rtio 10 r
LOUt5IA>N^ GULF CLUSTER MODEL CA.UBBATION
(MO/coT
(JOMM
Figure F2-2
Source: Reference 3
-------
I-1
TEXAS GULF CLUSTER MOOBL CM.IBRAT.OM
Source: Reference 3
Figure F2-3
-------
Ln
N)
S.79
TTOM2* I
ioso*«* I
NSt-wi\
I TR*.MS,Fr.O.
7.31 tUMlMlw.%uir\m
7.01 aox £Mimo*<%
(,TOH»)
SMALL M1DCOK1TINENT GUJ5TER MODEL OU&RAJION
Figure F2-4
Source: Reference 3
-------
Ul
LO
Pv»£W«i> NMUHM. uy
T?i 1
C« WO
wJr^»i° IP**
rissin-*,«,..c0.-
ll.«7
C« TO 200'F (..<
[ *~ ' .« FfWM 5)
MC4 *.OS
I6*6O Pl\C.MiUM
.11 UMUXOBO
.94 BT^
fro**.)
LAR&E MID^EbT CLUSTER MODEL CALIE.RATIOM
Figure F2-5
Source: Reference 3
-------
i t.4& *•* ro '«»'?
6.*7
3.99_ LP&
RtFIMVXV CAJt. TO
.CO PuPCWKfcCO iCx FO^ AvLKYLJ^TIOM
. n Pu^^CH^^to wic^
155.1* tltt-ICNTM.
1.94
F^OM TRANSFtB,
[V^E5TCO\bT CLUSTER MODEL CALIBRATION
| (MO/CO)
Source: Reference 3
Figure F2-6
-------
fe.79 ^PG_
BT *»•• '.".f-Ws^*!
pr tK-)(i.*M,A iO»
a.Tt UK!**'
4S.t>a i
5.*»ei p-trowKiEix re£O FHOM Tt^Avi^PGU.
II. lO CM". FfilO CROM
EK^T COK^T CLUSTER MODEL CALIBRATION
Figure F2-7
Source: Reference 3
-------
A. The cluster models are based on individual
existing and operating refineries. Thus,
assumptions and conclusions based on work
done with the models could be verified with
existing data or through testing in these
refineries.
B. An extensive data base from these refineries
and their operation already exists, and much
of it has been gathered and evaluated by EPA
and Arthur D. Little, Inc.
C. The methodology for calibrating these models
and scaling them up has been developed. The
procedures for altering or updating the models
are available.
D. The cluster models were developed in coopera-
tion with API and industrial representatives.
Thus, they are familiar with this simulation
method.
E. From the results of scaling-up these models,
it appears that they are accurate in simu-
lating the U.S. refining industry.
F. Eighteen operating refineries were used in
developing the cluster models. Eleven dif-
ferent companies are represented. Thus, a
significant pool of existing refineries is
identified as potential sampling sites.
56
-------
G. The refinery cluster models are all from dif-
ferent geographical areas of the country, but
each represents a site of significant refinery
operations. Environmental assessments based
on the cluster models could be developed under
different, yet realistic, atmospheric condi-
tions.
H. Each of the cluster models has, as its feed,
different composite crude oils. Thus, the
effect of various crude oils on refinery
operation and emission characteristics could
be evaluated.
I. The changing characteristics of the crude
feedstock and the product slates have been
projected through 1985. Future operating and
emission characteristics can be evaluated if
desired.
J. The processing configurations of each of the
refinery cluster models is different. Thus,
the characteristics of many different processes
can be evaluated while using only the six
cluster models.
There are some deficiencies and omissions in the
cluster model simulation method.
A. There are no topping refineries (EPA Type A
refineries) represented in the clusters.
There are over 130 of these refineries in
the U.S. Most of them are small, however,
57
-------
and the total refining capacity of these re-
fineries amounts to only 10 percent of the
total U.S. capacity.
B. Small refineries of less than 50,000 barrels/
day capacity are not represented in the cluster
models. Over half of the refineries in the
U.S. fall in this category. They refine only
about one-quarter of the total U.S. refinery
crude input, however.
C. The models were developed in 1974-1975. If
they are to be used for a detailed and com-
prehensive environmental assessment of re-
fineries, recalibration of the models might
be required. This could be determined from
a comparison of current operations in the 18
reference refineries to those of 1977-1978.
D. Like the two models discussed previously, no
method is available to accurately characterize
the number of fugitive sources within the
cluster models.
2.4 THE DEVELOPMENT OF REFINERY PLOT PLANS
Powell, et al.,1* have developed for the EPA plot
plans for four refinery cases. In this work, representative
refinery process and auxiliary units were placed into modules.
These modules were located on refinery plot plans for the
purpose of generating hydrocarbon ambient dispersion models.
58
-------
2.4.1 - Description of Plot Plans
Four refinery cases were developed. These were:
• A small capacity existing refinery with a crude
unit charge rate of 50,000 barrels per day.
• An intermediate capacity existing refinery with
a crude unit charge rate of 200,000 barrels per
day.
• A large capacity existing refinery with a crude
unit charge rate of 350,000 barrels per day.
• A relatively new refinery with a crude unit
charge rate of 250,000 barrels per day.
The following information was given for each of the
four refinery cases:
• A plot plan detailing the location and area
of each refinery module.
• The estimated number and description of all
potential emission sources within each module.
Included among these potential sources are
valves, fittings, pumps, compressors, heaters,
tanks, wastewater treating units, cooling
towers, and flares.
• Descriptions of tall and wide structures which
could affect airflow and, consequently, dis-
persion of pollutants.
59
-------
A summary of the type and number of potential emis-
sion sources is given in Table F2-14 for the four refinery cases,
The number of sources was estimated using several methods.
The number of pumps and compressors (both active and
spare) was determined by counting these sources on process flow
diagrams. Actual flow diagrams and generalized process flow
diagrams from published literature were both used. Where not
specified, the number of spare pumps and compressors was esti-
mated using engineering judgment.
The number of valves and fittings was estimated from
the number of pumps handling hydrocarbons in a process module.
The number of process valves was assumed to be 70 times the
number of active and spare pumps. The number of control valves
was determined from process flow diagrams or was estimated from
the number of control valves in similar processes.
The total number of screwed and flanged fittings was
derived by multiplying the number of active and spare pumps
(handling hydrocarbons) by 235. . The number of vessels and
columns within a module was used to estimate the number of
relief valves.
Other equipment and process information given for
each refinery module include: the number of process heaters
and boilers, the design heat input for each device, stack
heights and diameters, storage tank information, and the area
and dimensions of each process module.
The plot plan for the large existing refinery case
is shown in Figure F2-8. The individual modules are identified
in Table F2-15. A product slate for the large existing re-
finery case is presented in Table F2-16.
60
-------
TABLE "F2-14. TYPE AND NUMBER OF POTENTIAL EMISSION SOURCES
FOR FOUR REFINERY CASES
Refinery Type
Pumps - active and
spare in hydro-
carbon service
Compressors -active
and spare in
hydrocarbon
service
Pumps and com-
pressors - active
and spare in
aqueous service
Control valves
Process valves
Fittings
Relief valves
Sample valves
Combustion devices
Tanks
Flares
Small
Existing
Refinery
50,000
186
17
30
228
16,645
54,706
237
323
29
78
2
Intermediate
Existing
Refinery
Capacity,
200,000
411
37
43
364
28,240
94,867
429
634
49
416
4
Large
Existing
Refinery
bbl/day
350,000
614
61
53
809
48,670
162,996
739
997
91
236
3
New
Refinery
250,000
315
24
25
352
24,805
83,337
371
495
32
70
2
Source\~ Reference 4
61
-------
Capacity =350,000 bbl/day Area = 6120 ft x 9100 ft
; . 22i
<5
28 29 30 31
25 26 27
47
49
58
ss
50
51
61
52
52
35
36
37
39
36
4O
41 42
69
70
71
72
65
67
73
•n
14
15
12
16
17
18
19
20
21
55
56
57
75
76
Figure F2-8. Arrangement of refinery process and auxiliary
modules for a large existing refinery.
Source:! Reference A
62
-------
TABLE F2-15. DESCRIPTION AND IDENTIFICATION NUMBER OF MODULES
IN A LARGE EXISTING REFINERY
Module No. Description
LI Buffer Zone
L2 Feedstock Storage
L3 Crude Oil Storage
L4 Feedstock Storage
L5 Feedstock Storage
L6 Crude Oil Storage
L7 Feedstock and Product Storage
L8 Crude, Feedstock, and Product Storage
L9 Crude, Feedstock, and Product Storage
L10 Oil-Water Separator
Lll Product Storage
L12 Product Storage
L13 Distillation and Gas Recovery Unit
L14 Jet Hydrofiner/Catalytic Reformer
LI 5 Naphtha Nydrotreater
L16 Hydrotreater (It cycle oil)
L17 Hydrogen Manufacturing
L18 Partial Oxidation Unit
L19 Future Expansion
L20 Cooling Tower
L21 Flares
L22 Feedstock and Product Storage
L23 Naphtha Hydrotreater
L24 Vacuum Gas Oil Unit
L25 Benzene Fractionation
L26 Steam Rerun Stills
L27 Future Expansion
L28 Crude Distillation
L29, Catalytic Reformer
L30 Vacuum Residuum Desulfurizer
L31 Hydrogen Manufacturing
L32 Alkylation
L33 Distillate Hydrodesulfurization
(hvy gas oil)
L34 Sulfur Recovery
L35 Tanks/Cooling Towers
L36 Catalytic Reformer
L37 Aromatics Extraction
L38 Catalytic Cracking
L39 Para-Xylene Plant
LAO Delayed Coker
L41 Barrel Storage
- Continued
63
-------
Table F2-15. Continued
Module No
Description
L42
L43
L44
L45
L46
L47
L48
L49
L50
LSI
L52
L53
L54
L55
L56
L57
L58
L59
L60
L61
L62
L63
L64
L65
L66
L67
L68
L69
L70
L71
L72
L73
L74
L75
L76
Barrel Reconditioning
Feedstock Storage
Storm Water Impound Basin
Warehouse
Gas Holder/Blowdown Stack
Gas Holder/Blowdown Slack
Fire Prevention Training Facility
Oil-Water Separator
Asphalt Plant
Solvent Treating Plant/Boiler House
SO2 Treating Plant/Tanks
Lube Oil Packaging
Coke Storage
Crude Oil Storage
Feedstock Storage
Tanks/Impound Basin
Administration
Oil-Water Separator
Gasoline Sweetener/Crude Distillation
Crude Distillation/Crude Desalter
Specialty Crude Distillation
Specialty Crude Distillation/Condenser Box
Gasoline Fractionating Unit
Tank Loading/Truck Loading/Vapor Recovery
Buildings
LPG Storage and Blending
Vapor Recovery/Gasoline Rectifier/Tanks
Main Pump House
Product Storage
Wastewater Treatment
Building
Product Storage
Shops and Warehouse
Crude Oil Storage
Crude, Feedstock, and Product Storage
Source: Reference 4
64
-------
TABLE F2-16. PRODUCT SLATE FOR LARGE CAPACITY
EXISTING REFINERY
Crude Charge Rate
Benzene into Refinery
Product Slate
Product
Gasoline
Fuel oil
Jet fuel
Distillate fuel oil
Coke
Naphtha
Toluene and xylene
Asphalt
Mixed olefins
Benzene
Lube oils
Kerosene
350,000 bbl/day
1,220 bbl/day
Production Rate bbl/day
152,500
68,770
44,780
49,870
' 21,550
8,579
7,189
4,541
2,906
1,572
799
403
363,500
Source: Reference 4
65
-------
2.4.2 , Utility of the Refinery Plot Plans
The Refinery Plot Plans described in this section
have both advantages and disadvantages for potential applica-
tion in the Refinery Assessment Program. The advantages
include:
• Detailed refinery Layouts are presented.
These layouts can be used in association
with atmospheric dispersion models to
evaluate hydrocarbon and other pollutant
concentrations in the vicinity of refineries.
• Each refinery case contains a very detailed
listing of all refinery modules including
process units, auxiliary/supportive units,
and storage facilities.
• Detailed source counts are given for
each of the modules. This allows hydro-
carbon emissions to be estimated from
each refinery module.
• Heaters and boilers, along with stack
dimensions, are described in detail.
Rated heat duties are given for each
combustion source.
• Process streams through each pump and
compressor are classified into one of
three groups according to their volatility.
This is helpful in applying valve, puanp seal,
and compressor seal emission factors which
are functions of stream volatilities.
66
-------
1 Some disadvantages of the refinery plot plan cases
are :
• The numbers estimated for each source type
were not based on actual field counts.
• All module shapes and refinery layout are
rectangular.
• Standard size (capacity) modules were used
to develop the plot plans for the four re-
finery cases.
• The four refinery cases do not necessarily
correspond to any "typical" or "representa-
tive" set of model refineries.
2.5 THE PRODUCTION OF ORGANIC CHEMICALS AT REFINERIES
Processing modules for the production of organic
chemicals other than BTX (benzene/toluene/xylene), LPG, and
C3/Ci, olefins are not included in either the ADL cluster
models3 or the refinery cases developed by Powell, et al."
The capacities and sites of chemical production facilities were
surveyed to determine the need for the inclusion of additional
process modules in the existing refinery simulation models.
The 20 organic chemicals produced in the greatest
volume in the United States (1974) were identified. It was
found that 15 of these chemicals are produced in significant
quantities at refinery associated sites, often at facilities
operated by chemical subsidiaries of the parent refining com-
panies:. Production site and capacity summaries for these 15
67
-------
chemicals are tabulated in Table F2-17. The chemical manu-
facturing capacity at refinery associated sites exceeds 50
percent of the total domestic capacity for only five organic
chemicals: benzene, cuinene, propylene, toluene, and p-xylene.
In addition, it is probable that over half of the mixed
xylenes production capacity is also located at refinery as-
sociated sites, but individual site capacities were not avail-
able to verify this assumption. Only benzene, toluene and
propylene are produced at more than 20 refinery associated
sites.
Chemical production at sites associated with Types C
(petrochemical) and E (integrated) refineries was identified.
Data on production capacities for organic chemicals at these
locations were obtained from the 1976 Directory of Chemical
Producers.7 A total of 30 refinery sites were considered. Only
seven organic chemicals were produced at more than 6 of the 30
refinery associated locations. These chemicals were benzene *
toluene, mixed xylenes, propylene, cyclohexane, cumene and pro-
pane. The refinery associated production sites are listed in
Table F2-18. Only benzene, toluene, mixed xylenes and
propylene are produced at more than 10 refinery associated
sites.
With the exception of BTX and propylene production,
then, it is difficult to characterize the manufacture of
organic chemicals at refinery associated sites. The production
modules for aromatics and C3/C<. olefins are already included in
the refinery models. In both cases, the inclusion of additional
fractionation in the existing modules would provide the means
for producing benzene, toluene, mixed xylenes and propylene.
The addition of other chemical production modules to the exist-
ing refinery models does not appear to be justified.
68
-------
TABLE F2-17. CHEMICAL PRODUCTION AT REFINERY ASSOCIATED SITES
Chemical
Benzene
Butadiene
Cumcne
Ethylbenzene
Ethylene
Ethylene
Uichloride
Ethylene Glycol
Ethylene Oxide
Phenol
Propylene
Styrene
Toluene
Vinyl Chloride
Mixed Xylenes
P-xylene
Total 1976
Production
Capacity
(Domestic)
1,475 MM gal
4,406 MM Ib
3,785 MM Ib
8,406 MM Ib
25,220 MM Ib
12,880 MM Ib
3,400 MM Ib
3,960 MM Ib
2,556 MM lh
13,223 MM Ib
7,075 MM Ib
967 MM gal
6,165 MM Ib
(6,060 MM Ib)*
2,530 MM Ib
Production Capacity at
Refinery Associated Sites
Amount
1,099 MM gal
700 MM Ib
1,995 MM Ib
1,246 MM Ib
8,075 MM Ib
2,435 MM Ib
265 MM Ib
475 MM lh
238 MM Ib
9,031 MM Ib
1,010 MM Ib
833 MM Ib
1,540 MM Ib
2,230 MM Ib
% of
Total
Capacity
75
16
53
15
32
19
8
12
9
68
14
86
25
88
Total
No. of
Production
Sites
52
23
13
17
34
17
16
14
17
60
13
47
14
32
11
Refinery Associated
Production Sites
Number
35
4
9
6
10
3
1
1
3
37
3
27
2
18
10
% of
Total
Sites
67
17
69
35
29
17
6
7
18
62
23
57
14
56
91
1974 Production
Source: References 7, 12
-------
TABLE F2-L8. ORGANIC CHEMICALS PRODUCED AT REFINERY ASSOCIATED SITES
Refineries /Sites
•i'
Standard Oil Co. (California)
El Segundo, California
Getty Oil Company
Delaware City, Delaware
Amoco Oil Company
Wood River, Illinois
Clark Oil & Refining Company
Blue Island, Illinois
Skelly Oil Company
El Dorado, Kansas
Ashland Petroleum Company
Tonowanda, New York
Standard Oil Co. (Kentucky)
Pascagoula, Mississippi
Exxon Company
Linden, New Jersey
Texaco, Inc.
Westvil.le, New Jersey
Ashland Petroleum
Cattletsburg, Kentucky
Sun Oil Co. of Pennsylvania
Toledo, Ohio
Mixed
Benzene Toluene Xylencs Propylene Cyclohexane Cumene Propane
t i-
P/C P C C
p
C
C C
p p p p p
C C C C
C
p p p p
C C C C C
p p
Continued
-------
TABLE F2-18. Continued
Refineries/Sites
Mixed
Benzene Toluene Xylenes Propylene Cyclohexane Cumene Propape
Amoco Oil Company
Texas City, Texas
Coastal States Petrochemical Co.
Corpus Christi, Texas
Cosden Oil & Chemical Company
Big Spring, Texas
Marathon Oil Company
Texas City, Texas
Phillips Petroleum Company
Borger, Texas
Phillips Petroleum Company
Sweeny, Texas
Suntide Refining Company
Corpus Christi, Texas
Amoco Oil Company
Yorktown, Virginia
Atlantic Richfield Company
Carson, California
Shell Oil Company
Martinez, California
Standard Oil Co. (California)
Richmond, California
Continued
-------
TABLE F2-L8. Continued
Refineries/Sites Benzene Toluene Xylenes PropylenR Cyc.lohexane Curaene Propane
Phillips Petroleum Company
Kansas City, Kansas
Exxon Company
Baton Rouge, Louisiana
Sun Oil Company
Marcus Hook, Pennsylvania
Atlantic Richfield Company
Houston, Texas
Exxon Company
Baytown, Texas
Gulf Oil Company
Port Arthur, Texas
Union Oil Company
Nederland, Texas
P = Produced by petroleum company.
C = Produced by chemical subsidiary of petroleum company.
-------
2.6 SUMMARY
The cluster model refineries developed by ADL3 and the
refinery cases developed by Powell, et al. , "* seem best suited
for emissions evaluation and an environmental assessment. The
cluster models are useful for preliminary emissions evaluation
to determine the relative importance of various types of
emission sources.
The refinery plot plans and refinery cases developed
by Powell, et al.,1' are quite detailed. They were developed
for use with atmospheric dispersion models. Consequently,
they should be suited for use in an environmental assessment.
Actual field source counts have been obtained as a
part of this program and are provided in Section 4 for a
variety of process units. When these counts are compared to
those developed by Powell, et al.,* it is evident that there
are considerable differences between the two estimates. This
is probably a result of real differences between the actual
number of sources in process units of the same type as well as
differences resulting from the methods used to obtain these
estimates.
Estimates for total emissions from a variety of pro-
cess units are listed in Section 4 of this Appendix. Where
possible, both sets of estimates for fugitive emission sources
are included. It is difficult to say which estimate is the
more representative of the two. Since the results of this pro-
gram indicate that there are wide variations in the number of
sources between similar process units in different refineries
and that these differences do not correlate with the unit
capacity, the use of an "average" number of sources per process
73-
-------
unit is probably not appropriate. Therefore, accurate esti-
mates of total fugitive emissions from actual process units
should be obtained using the number and distribution of sources
actually present in that particular process unit.
2.7 REFERENCES FOR SECTION 2
1. Cantrell, Ailleen. Annual Refining Survey. Oil and Gas J.,
77(13): 122-156, March 26, 1979.
2. Halper, Martin. Development Document for Effluent Limita-
tions Guidelines and New Source Performance Standards for
the Petroleum Refining Point Source Category. PB 238 612,
EPA Contract No. 68-01-0598. Environmental Protection
Agency, Washington, B.C., April 1974.
3. Arthur D. Little, Inc. Impact of SOX Emissions Control on
Petroleum Refining Industry. Final Report, Vol. II, De-
tailed Study Results. EPA-600/2-76-161b. Cambridge,
Massachusetts, June 1976.
4. Powell, D., et al. Development of Petroleum Refinery Plot
Plans. EPA-450/3-78-025. Pacific Environmental Services,
June 1978.
5. Bombaugh, K. J., et al. Sampling and Analytical Strategies
for Compounds in Petroleum Refinery Streams. Final Report,
Volumes, PB 251 744 (Vol. 1); PB 251 (Vol. 2); EPA-600/2-
76-012, Radian Corporation, Austin, Texas, January 1976.
6. U.S. Bureau of Mines. Crude Petroleum, Petroleum Products,
and Natural Gas Liquids, December 1974 Mineral Industry
Surveys. Petroleum Statement Monthly, Washington, D.C.,
April 1975.
7. Stanford Research Institute. 1976 Directory of Chemical
Producers, U.S.A. Menlo Park, California, 1976.
8. Arthur D. Little, Inc. The Impact of Producing Low-Sulfur,
Unleaded Motor Gasoline on the Petroleum Refining Industry.
Final Report, Vol. II, Detailed Study Results. EPA-450/
3-76-015B. Cambridge, Massachusetts, May 1976.
9. Arthur D. Little, Inc. The Impact of Lead Additive Regula-
tions on the Petroleum Refining Industry. Final Report,
Vpl. II, Detailed Study Results. EPA-450/3-76-016b.
Cambridge, Massachusetts, May 1976.
74
-------
10. Cahtrell, Ailleen. Annual Refining Survey. Oil and Gas J. ,
71(14): 99-125, April 2, 1973.
t
11. Cantrell, Ailleen. Annual Refining Survey. Oil and Gas J.,
72(13): 82-106, April 1, 1974.
12. Cantrell, Ailleen. Annual Refining Survey. Oil and Gas J. ,
74(13): 124-156, March 29, 1976.
75
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SECTION 3
REFINERY RAW MATERIALS AND PRODUCTS
Each refinery has its own set of raw materials and
products which varies with availability and market demand. The
general categories and characteristics of each will be described
in this section. A detailed description is beyond the scope of
this report.
3.1 CHARACTERISTICS OF RAW MATERIALS
The raw materials for a refinery include crude oil,
catalysts, acids, bases, light olefins, and gasoline additives.
All are used in varying amounts.
3.1.1 Crude Oil
Crude oil is the basic raw material of any refinery.
Its definition, according to ASTM D-288, is as follows:
"A naturally occurring mixture, consisting
predominantly of hydrocarbons, and/or of
sulfur, nitrogen and/or oxygen derivatives
of hydrocarbons, which is removed from the
earth in liquid state or is capable of being
removed.
Crude petroleum is commonly accompanied by
varying quantities of extraneous substances
such as water, inorganic matter and gas.
The removal of such extraneous substances
alone does not change the status of the
mixture as crude petroleum. If such removal
76
-------
appreciably affects the composition of the
oil mixture, then the resulting product is
no longer crude petroleum."
Crude is primarily a liquid; it may be light and easy-
flowing or a heavy, semi-solid substance. Its color may be
green, brown, black, or almost white. Specific gravities of
crudes usually range between 0.8 and 0.95; most kinematic
viscosities are between 0.023 and 0.23 stoke.1
uniform.
The elemental composition of crude is amazingly
The ultimate analysis is given in Table F3-1.
TABLE F3-1.
ULTIMATE ANALYSIS OF CRUDE OILS,
WEIGHT PERCENT
Carbon
Hydrogen
Sulphur
Nitrogen
Oxygen
Metals (Fe, V, Ni,
83.9
11.4
0.06
0.11
0.5
etc.) 0.03
- 86.
- 14.
- 8.
- 1.
8
0
00
70
Source: Reference 2
3.1.1.1 Major Types--
There are as many different crudes as there are
sources. Though the elementary composition is relatively uni-
form, small variances can greatly affect physical properties.
Because crudes are so very complex and varied, there is no one
simple^method for classification. It has been proposed that
77
-------
classification systems now in use be replaced by the concept of
a continuous spectrum of oils arranged in order of increasing
cyclization of the components.3 Two methods for categorization
are presented here: the classification by product commonly used
by refiners and a geographic classification.
Classification By Product--Refiners often classify
crudes according to the residue from their nondestructive dis-
tillation. This classification method usually gives some indi-
cation of the ultimate products. There may be as many as nine
categories in the classification to account for the various
types of oils. Some oils are combinations of several types.
The simplest classification has two categories:
paraffin-base and asphalt-base. Paraffinic crudes are good
sources of paraffin wax (condensed paraffins) and high-quality
motor oils and kerosenes. Asphaltic crudes are good sources of
asphalt (primarily condensed aromatics), machine lubricating
oils and high-quality gasoline. Asphaltic crudes usually have
high sulfur, oxygen, and nitrogen content; light and inter-
mediate fractions of asphaltic crude are higher in naphthenes.
Geographic Classification--Crudes from a certain geo-
graphic region usually have similar characteristics. Therefore,
it is often convenient to classify crudes according to their
region of origin.
No one geographic classification is used universally.
However, the classification used by A. D. Little5 is considered
the most widely used one. This classification is presented as
Table F3-2. Some general properties of the crudes categorized
in Table F3-2 are listed in Table F3-3.
78
-------
TABLE F3-2. GEOGRAPHIC CLASSIFICATION OF CRUDES
Name
Type
Louisiana
Oklahoma
West Texas
Wilmington
Ventura
Nigeria Forcados
Algerian Hassi
Messaoud
Arabian Light
Minas
Tia Juana Medium
Louisiana and low-sulfur Texas crudes
Light, sweet crudes from the mid-
continent
High-sulfur crudes from Texas and
New Mexico
Heavy crudes from California
Light crudes from California
Heavy African crudes
Light African crudes
Considered representative of average
Middle East crudes
Considered representative of South-
east Asian crude imports
Considered representative of
Venzuelan crude exports
Source: Reference 5
79
-------
TABLE F3-3.
GENERAL PROPERTIES OF CRUDES ACCORDING
TO GEOGRAPHIC LOCATIONS
00
o
Domestic Crudes
Yield, Volume %
Methane/ethane (FOE)
Propane
Isobutane
Normal butane
Straight-run naphtha:
C5-160°F
Light 160-2008F
Medium 200-340°F
Heavy 340-375°F
Light gas oil 375-500°F
Heavy gas oil 500-650°F
Vacuum overhead 650-1050°F
Vacuum bottom 1050°F+
Crude properties
Gravity (°API)
% weight sulfur
Louisiana
—
0.20
0.40
0.70
4.09
2.99
13.05
3.47
17.50
19.50
32.50
5.60
36.2
0.2a
West
Texas
Sour
0.04
0.50
0.40
1.30
5.70
3.80
15.50
3.30
13.39
14.11
29.60
12.30
33.4
1.31a
Oklahoma
—
0.48
0.46
2.18
7.70
4.76
16.66
3.93
13.76
11.77
28.04
10.25
40.2
0.212
California
Wilmington
0.002
0.093
0.095
0.318
2.09
2.47
7.64
1.50
9.29
11.96
38.54
26.00
19.6
1.28
California
Ventura
0.005
0.400
0.295
1.069
5.72
3.25
14.26
3.34
11.54
12.53
32.08
15.50
29.7
1.56
Alaskan
North
Slope
0.06
0.28
0.13
0.49
3.53
2.57
9.25
2.69
12.43
15.50
29.49
23.52
27.5
0.96
(Continued)
-------
TABLE F3-3. Continued
CO
Foreign Crudes
•V'
Yield, Volume %
Methane/ethane (FOE)
Propane
Isobutane
Normal butane
Straight-run naphtha:
C5-160°F
Light 160-200°F
Medium 200-340°F
Heavy 340-375°F
Light gas oil 375-500°F
Heavy gas oil 500-650°F
Vacuum overhead 650-1150°F
Vacuum bottom 1050° F-f
Crude properties
Gravity (°API)
% weight sulfur
Nigerian
Forcados
0.04
0.04
0.51
0.79
2.70
3.40
11.70
2.80
18.10
20.60
30.40
8.50
29.4
0.21
Arabian
Light
—
0.17
0.17
1.06
4.85
3.27
14.93
3.95
13.39
15.01
29.50
13.70
34.6
1.7
Venezuelan
Tia Juana
0.01
0.59
0.27
0.45
3.89
2.20
9.09
2.52
10.30
12.70
32.80
25.00
26.3
1.51
Algerian
Hassi
Messaoud
0.04
1.21
0.53
3.27
8.29
5.00
21.19
5.02
15.78
11.92
22.71
4.99
44.7
0.13
Mixed
Canadian
0.05
1.13
0.49
1.98
6.60
3.89
16.50
3.51
H.40
15.20
25.60
10.60
39.0
0.55
Tndo- i , ,•
nesian Natural
Minas Gasoline
0.003
0.139
0.141 4.71
0.379 7.58
1.70 62.02
1.60 13.99
9.10 11.70
2.60 —
10.70
15.00
41.00
18.00
35.3
0.07
Radian calculation.
Source: Reference 4
-------
}" Crudes are also often referred to by .specific field.
3.1.1.2 Major Constituents--
As stated previously, crudes are basically very com-
plex mixtures of hydrocarbons. There are probably more than
3000 compounds in any one crude.6 Because the composition of
crude influences so many aspects of refinery design and opera-
tion, it has been the subject of much research in recent years.
The hydrocarbons in the crude are normally paraffins,
naphthenes (cycloparaffins) and aromatics. Their structures
may be simple or complex. There is evidence of normal paraffins
with as many as 78 carbon atoms. The concentration of branched
paraffins appears to decrease with increasing molecular weight.
There also appears to be a tendency toward methyl branches at
the ends of molecules.7
Naphthenes usually have 5-, 6-, or 7-membered rings,
whether they are monocyclic or polycyclic. Methyl substitution
is quite common. Most of the alkyl derivatives of benzene,
naphthalene, biphenyl, and phenanthrene with boiling points up
to 300°C have been found in crude. Substituents on the ring
usually consist largely of methyl groups with one long, rela-
tively unbranched, alkyl group. In high boiling point fractions,
aromatic rings and cycloalkyl rings may be combined.7
The hydrocarbons in crude are usually categorized
according to their boiling point ranges. The terminology used
to describe the fractions varies slightly. Exact boiling point
cut-offs vary with the intended use for the fraction. Two
example slates of fractions are given in Tables F3-4 and F3-5.
82
-------
TABLE F3-4. BOILING RANGES OF TYPICAL CRUDE OIL FRACTIONS
Fraction Boiling Range, °F
Light gases -259 to -44
(methane, ethane, some propane)
Propane -44
Butanes 11 to 31
Light Naphtha 30 to 300
Heavy Naphtha 300 to 400
Kerosine 400 to 500
Light Gas Oil 400 to 600
Heavy Gas Oil 600 to 800
Vacuum Gas Oils 800 to 1100
Residue 1100+
(source of asphalts or waxes)
Source: Reference 1
TABLE F3-5. BOILING RANGES OF TYPICAL CRUDE OIL FRACTIONS
Fraction Boiling Range, °F
Butanes and lighter Not given
Light straight-run gasoline (LSR) 90 to 190
Naphtha (heavy straight-run gasoline) 190 to 380
Kerosine 380 to 520
Light gas oil (LGO) 520 to 610
Atmospheric gas oil 610 to 800
Vacuum gas oil (VGO) 800 to 1,050
Vacuum reduced crude 1050+
Source: Reference 8
83
-------
The hydrocarbon composition of a crude fraction
depends on the type of crude. This is illustrated in Table
F3-6, which shows the approximate hydrocarbon compositions of
some representative fractions in paraffinic and asphaltic
crudes.
3.1.1.3 Potentially Hazardous Constituents--
The potentially hazardous constituents of crude oil
are of four types: hydrocarbons, sulfur compounds, nitrogen
compounds, and trace elements. Unlike coal, oil is subjected
to processing during which essentially all components are con-
verted into usable products. The hazardous constituents may be
released during any of the processing steps. Since most oil
products are fuels, the hazardous constituents may also be
released when these fuels are combusted.
Hydrocarbons--Table F3-7 is a list of hydrocarbons
which have been identified in crude and which have been desig-
nated as potentially hazardous. The compounds included have
either been assigned a threshold limit value (TLV) by the
American Conference of Governmental Industrial Hygienists or
assigned a rating of 2 or 3 (capable of causing permanent
damage to humans) by Irving Sax in Dangerous Properties of
Industrial Materials, 1975 edition.
Sulfur Compounds--Organic sulfur compounds are present
to some extent in all crudes, but the amount can vary from 0.06
to over 8 weight percent. The sulfur compounds are not dis-
tributed evenly through the crude: in general, the concentra-
tion of sulfur compounds increases as molecular weight increases
The complexity of sulfur-containing molecules also increases
with increasing molecular weight.2
84
-------
TABLE F3-6.
DISTRIBUTION OF HYDROCARBON TYPES
IN SOME PETROLEUM FRACTIONS
Boiling Range
502 ASTM Paraffin-Base Crude,
Dlstlllntlon, Ut 7.
Fraction
Gasoline
Kerofllne
Cas oil
Heavy
Distillate
°F Paraffins
280 65
450 60
600 35
750 20
Naphthenes
30
30
55
65
Aromatics
5
10
15
15
Asphalt-Base Crude,
Wt 7.
Paraffins Naphthenes
35 55
25 50
65
55
Aromatics Unsaturated
10
25
33 2
43 2
OO
L/l
Source: Reference 1
-------
TABLE F3-7. POTENTIALLY HAZARDOUS3 HYDROCARBONS
IN CRUDE OIL
Compound Concentration
Methane T
Ethane T
Propane m
Methylpropane T
Butane m -»• M
Methylbutane m
n-Pentane tn
2 , 2-Dimethylpropane T •* m
n-Hexane m
2-Methylpentane m •* M
3-Methylpentane m
2,2-Dimethylbutane m
2,3-Dimethylbutane m
n-Heptane m-M
2,3-Dimethylpentane m
2,4-Dimethylpentane m
n-Octane m
2-methylheptane m
2,3-Dimethylhexane m
2,4-Dimethyhexane m
2,2,4-Trimethylpentane T
n-Dodecane T •* m
Cyclopentane m
Cyclohexane m
Methylcyclohexane m -»• M
Cycloheptane m
Benzene T -»• m
Toluene T -»• m
Ethylbenzene T •* m
Dimethylbenzene (Xylene) T •* m
n-Propylbenzene m
Isopropylbenzene (Cumene) m
1,2,3-trimethylbenzene m
1,2,4-Trimethylbenzene m
1,3,5-Trimethylbenzene m
Isobutylbenzene m
sec-Butylbenzene m
tert-Butylbenzene m
l-Methyl-5-Isopropylbenzene m
1,2-Diethylbenzene m
1,3-Diethylbenzene m
(Continued)
86
-------
TABLE F3-7. Continued
Compound
Concentration
1,4-Diethylbenzene
l-Methyl-4-tert-butylbenzene
1-MethyInaphthalene
2-Methylnaphthalene
Pyrene
Coronene
Benzo(e)pyrene
1,2,3,4-Tetrahydronaphthalene
Biphenyl
Acenaphthene
Benzofluorenes
Phenanthrene
Benzophenanthrene
Naphthenophenanthrenes
Dinaphthenophenanthrenes
Trinaphthenophenanthrenes
Tetranaphthenophenanthrenes
Pentanaphthenophenanthrenes
Fluoranthrenes
Perylene
Phenyleneperylene
Dibenzoperylene
Chrysene
Benzo(g)chrysene
3-Methylchrysene
Naphthenochrysenes
Anthracene
Benzanthracene
m
T
T
T
T
T
T
T
T
T
T
T
T
T
T
T
T
T
T
T
T
T
T
T
T
T
T
T
Sources: References 3, 9, 10, 11, 12, 13, 14, 15, 16
aThe compounds included in this list have either been
assigned a Threshold Limit Value (TLV) by the American
Conference of Governmental Industrial Hygienists or
assigned a rating of 2 or 3 (capable of causing permanent
damage) by Irving Sax in Dangerous Properties of Indus-
trial Materials, 1975 edition.
Key to Concentrations:
T = trace: <100 ppm
m = minor: 100 ppm to 2.99%
M = ma j or : >_3 . 0%
87
-------
Trace quantities of 48 thiols, almost 200 sulfides,
and a number of sulfites, sulfonates, and sulfones have been
identified in crude and identified as potentially hazardous ac-
cording to footnote "a" in Table F3-7. Hydrogen sulfide is known
to be a dangerous chemical. Many of the sulfur compounds in
crude are decomposed by normal refinery processes.
Hydrogen sulfide is often associated with the lighter
gases. Elementary sulfur can also be dissolved in the oil,
possibly formed by the oxidation of hydrogen sulfide.
In lower boiling point fractions (up to about 400°F),
mercaptans (thiols) appear to predominate. Alkyl mercaptans
with the thiol group on a primary carbon are most prevalent in
fractions with boiling points up to 200°F; in higher boiling
fractions, the thiol group is more likely attached to a second-
ary carbon.
Cyclic mercaptans appear in the kerosine range; thio-
ethers and cyclic sulfides in naphthenes. In higher boiling
fractions, the tendency appears to be toward sulfur substitution
in saturated rings. Little is known of the sulfur species in
the heaviest distillates; they are believed to be mostly poly-
cyclic molecules with sulfur substitutions in the rings.
Nitrogen Compounds—As with sulfur compounds, the
concentration of nitrogen compounds in crude increases with
increasing boiling point. Unlike sulfur compounds, nitrogen
compounds do not decompose in normal refinery processes. Also,
nitrogen content is lower than sulfur content — less than one
percent.
88
-------
i Approximately one-fourth to one-third of the nitrogen
in crude is found in basic nitrogen compounds. The basic
nitrogen compounds which have been isolated, mainly in gas oils,
are alkyl, mainly methyl or dimethyl, substituted quinolines and
pyridines. All of these alkyl quinolines and some of the pyri-
dines have been designated potentially hazardous as defined by
footnote "a" in Table F3-7. The nonbasic compounds indole and
the carbazoles have also been identified in crude oils and have
been designated as potentially hazardous.
Oxygen Components—Crude oil generally contains less
than 2 percent oxygen. However, oxygen content increases with
increasing molecular weight: residual oils may contain as much
as 8 percent oxygen.2
The oxygen compounds in crude designated as poten-
tially hazardous include the lower molecular weight carboxylic
acids and alkyl ketones, some cyclic ketones and phenols.
These compounds are present in trace amounts in crude.
Trace Metals--Trace quantities, most less than 100
ppm, of a number of metals have been found in crude. These
elements are usually observed in the ash remaining from combus-
tion. Twenty-eight metals which have been found by spectro-
graphic examination of ash are listed in Table F3-8. Of the
metals listed, vanadium, nickel and iron are usually present in
the greatest quantities. Table F3-9 gives relative trace
element concentrations in several representative crudes. Most
of the elements are considered potentially hazardous.
89
-------
TABLE F3-8.
TRACE METALS FOUND BY SPECTROGRAPHIC
ANALYSIS OF THE ASH FROM CRUDE OIL
Ag
Al
As
B
Ba
Ce
Co
Cr
Ca
Cu
Fe
Ga
K
La
Y-B
Mn
Mo
Na
Nd
Ni
Pb
Sn
Sr
Tl
V
Zr
Zn
U
Source: Reference 17
TABLE F3-9.
TRACE ELEMENT CONTENTS OF SOME CRUDE OILS
BY NEUTRON ACTIVATION ANALYSIS
Concentration
Element
V
Cl
I
Na
K
Mn
Cu
Ga
As
Br
Mo
Cr
Fe
Hg
Se
Sb
Ni
Co
Zn
Sc
u
California
(Tertiary)
7
1
-
13
-
1
0
0
0
0
-
0
68
23
0
0
98
13
9
0
-
.5
.47
-
.2
-
.20
.93
.30
.655
.29
-
.640
.9
.1
.364
.056
.4
.5
.76
.0088
-
in Crude Oil (ppmw)
Libya
8.
1.
--
13.
4.
0.
0.
0.
0.
1.
--
0.
4.
--
1.
0.
49.
0.
62.
0.
0.
2
81
0
93
79
19
01
077
33
0023
94
10
055
1
032
9
0003
015
Venezuela Alberta
(Boscan) (Cretaceous)
1100.
--
--
20.
--
0.
0.
--
0.
--
7.
0.
4.
0.
0.
0.
117.
0.
0.
0.
--
3
21
21
284
85
430
77
027
369
303
178
692
0044
0
25
1
2
-
0
-
-
0
0
-
-
0
0
0
-
0
0
0
-
-
.682
.5
.36
.92
-
.048
-
-
.0024
.072
-
-
.696
.084
.0094
-
.609
.0027
.670
-
-
Source: Reference 17
90
-------
3.1.2 '-.- Other Raw Materials
Besides crude oil, a large number of chemicals are
used at a refinery as treating agents, solvents, catalysts, and
additives. Many of these chemicals are eventually emitted to
the atmosphere.
3.1.2.1 Catalysts--
Petroleum refineries are the largest user of catalysts
in the chemical industry.1 The catalysts may be solid or
liquid. Table F3-10 lists the commonly used catalysts and the
processes in which they are used.
Catalyst fines are emitted to the atmosphere during
catalyst regeneration. Most catalysts are regenerated only a
few times a year; therefore, the escaping catalyst fines are
considered insignificant. Fluid catalytic cracking catalysts,
on the other hand, are regenerated continuously. In this case,
particle collection devices are used extensively to control the
emissions, both for environmental protection and for economic
reasons.
3.1.2.2 Gasoline Additives--
In many refineries, gasoline is the primary product.
Several chemicals are usually added to this gasoline to improve
its various qualities. These additives are described below:1'2
Antiknock Compounds--Tetraethy1 lead is the most
common antiknock compound; it was used in concentrations of up
to 0.08 volume percent before the lead phase-out program. By
1979, the maximum concentration approached 0.01 volume percent.
91
-------
TABLE F3-10. PRINCIPAL APPLICATIONS OF CATALYST MATERIALS3
to
»'
Catalyst Material
Aluminia
Aluminum chloride
Antimony trichloride
Bauxite
Bentonite clay
Clay
Cobalt -mo lybdena
Cobalt molybdate
Cobalt oxide
Copper
Copper pyrophosphate
Hydrochloric acid
Hydrofluoric acid
Iron oxide
Kaolin clay
Magnesia
Molybdena
Molybdenum
Nickel sulfide
Phosphoric acid
Platinum
Potassium
Silica-alumina
Sulfuric acid
Tungsten nickel sulfide
Crack-
ing
X
X
X
X
X
X
X
X
X
X
Processing
Reform- Hydro-
ing treating
X X
X
X X
X
X
X
X
X X
X X
X
Application
Isomeri- Alkyla- Polymer-
zation tion ization
X X
X
X
X
X
X X
X
X
X
X
aMany catalyst materials are also used for other purposes in a refinery.
Source: Reference 1
-------
Othersiwhich may be used are tetramethyl lead, triethylmethyl
lead, diethyldimethyl lead, and ethyltrimethyl lead. Ethylene
dibromide is added in a 1:1 relationship with the lead to cause
the formation of volatile lead bromide instead of lead oxide
which could foul the engine.
Antioxidants--Aromatic amines and alkyl-substituted
phenols are used in concentrations of about 50 ppm to inhibit
oxidation reactions which cause the formation of gums.
Metal Deactivators--Metal deactivators prevent the
dissolution of metals, particularly copper, by the gasoline or
deactivate metals already dissolved in the fuel. Copper in
particular catalyzes the formation of gums. Many metal deacti-
vators contain amine or diamine groups which combine with the
metal to form a compound which does not react further.
Anti-Corrosion Additives--Surface-active chemicals
are added to gasoline to emulsify any water present so that it
will not rust pipelines, storage tanks, and engine parts.
Antistall Additives--Antistall additives are used to
prevent carburetor icing. There are two types: cryoscopic
additives such as methanol, ethanol or isopropanol lower the
freezing point of condensed water so that it does not freeze;
surfactants such as polyalkylene glycols, alkyl phosphates, and
alkyl amines form a monomolecular hydrophobic film over metal
surfaces to prevent the adherence of ice crystals. Surfactants
are used at levels of 30 to 150 ppm; cryoscopic additives are
used at levels of about 0.05 to 1.0 percent volume.
93
-------
i.- Antipreignition Agents — Deposit modifiers such as
phosphorus compounds are used to change the character of
combustion-chamber deposits. These are used at about 200 pptn.
Lubricants--About 0.2 to 0.5 percent of a light
lubricating oil is often added to gasoline to lubricate the
intake valves and the top ring belt area of the engine.
3.1.2.3 Other Chemicals--
Many chemicals are used in the various refining pro-
cesses. Table F3-11 is a summary of major chemicals and their
principal uses. Some catalysts discussed above are included in
this table. The chemicals may be used alone or with other
chemicals.
3.2 CHARACTERISTICS OF REFINERY PRODUCTS
3.2.1 Final Products
According to an API study a few years ago, almost
3000 individual products are produced in petroleum refineries.1
Although there are many variations, these products are generally
grouped into a few major categories. These categories will be
described in this section.
3.2.1.1 Refinery Gases--
The most volatile fraction of crude oil processing is
a mixture of gases in the C: to C5 range. These gases may have
been a part of the original crude, they may have been injected
into the crude in the oilfield, or they may be a product of the
processing of heavier fractions. All are considered potentially
hazardous.2
94
-------
.TABLE F3-11.
MAJOR CHEMICALS USED IN REFINING AND
THEIR PRINCIPAL USES
Chemical
Uses
Acetic Acid
Acetone
Aluminum Chloride
Aluminum Oxide (Bauxite)
Aluminum Naphthenates
Aluminum Phenates
Aluminum Soaps
Aluminum Stearate
Barium Hydroxide
Break up emulsions
Increase treating of sulfuric acid
Reduce sulfur content
Extract polymers from cracked
distillates
Separate waxes
Regenerate clays
Isolate benzene in azeotropic
distillation
Solvent in determining oil content
of waxes
Cracking, alkylation, and iso-
merization catalyst
Cracking catalyst
Detergent additive for lubri-
cating oils
Treat spent caustic solutions
Neutralize acid-treated oils
Precipitate naphthenic acids
Prevent foaming before caustic
soda treating for mercaptan
removal
Remove inorganic salts from
furfural before refining
(Continued)
95
-------
TABLE F3-11. Continued
Chemical
Uses
Barium Salts
Benzene
Bore Char
Cadmium-Ammonium
Chloride
Cadmium Hydroxide
Cadmium Chloride
Cadmium Sulfide
Cadmium Oleate |
Cadmium Naphthenate )
Cadmium Dithiocarbamate
Cadmium Sulfonate
Calcium Oxide
Calcium Hydroxide
Calcium Carbonate
Calcium Chloride
Calcium Hypochlorite
Oxidation inhibitors, detergent
additives in lube oils
Solvent extraction to improve
viscosity index of lube oils
and remove waxes
Decolorize oil
Distillate Desulfurizing
Oxixation inhibitor in lube oil
Detergent additive
Neutralize acid-treated oils
Remove hydrogen sulfide and
organic acids from oils
Dessicant
Oxidize sulfides and mercaptans
in oils
(Continued)
96
-------
TABLE F3-11. Continued
Chemical
Uses
Chlorine
Clays
Cupric Chloride
Cresol
Dichloroethyl Ether
E t hano1amine s
(MEA, DEA, TEA)
Ethylene Bichloride
Ethylene Glycol
Formaldehyde
Oxidize disulfides to sulfonyl
halides and to remove mercaptans
Regenerate Bentonite clay
Regenerate sodium plumbite "doctor
solution"
Prepare calcium and sodium hydroxide
Improve cetane number of fuels
Adsorbents to improve color, odor,
and stability of waxes and lube
oils
Cracking catalysts
Convert mercaptans to insolubile
disulfides
Extraction of high-viscosity-
index, light-color, low-carbon-
residue lubricants from residual
or distillate base stock
Solvent in chlorex extraction to
improve viscosity index and
yields of paraffinic oils
Removal or recovery of water,
hydrogen sulfide, or carbon
dioxide from gaseous streams
Removing wax from lube oil
Selective recovery of benzene,
toluene, and xylenes from
petroleum shocks
Laboratory reagent and solvent
(Continued)
97
-------
TABLE F3-11. Continued
Chemical
Uses
Furfural
Hydrogen
Hydrogen Fluoride
Methyl Ethyl Ketone (MEK)
Methyl Isobutyl Ketone
(MIBK)
Natural Oils
Nitrobenzene
Phenol
Phosphorous Compounds
Phosphorous Pentoxide
Extraction of diesel fuels, burning
oils, cracking stocks, and crude
oils
Removal of low-cetane materials,
unstable and acidic materials,
sulfur, organometallic and
nitrogen compounds.
Extraction of aromatic, naphthene,
olefinic, and unstable hydro-
carbons from lube oils.
Hydrotreating
Hydrocracking
Hydroalkylation
Alkylation Catalyst
Remove wax from oils
Deoiling high-quality waxes
Production of lubes and greases
Extract carbon and sludge-forming
compounds from lube oils
Extraction of high-viscosity-
index, high-color, low-carbon-
residue lubricants from residual
or distillate base stock
Improve viscosity index, color and
oxidation resistance, and to
reduce carbon and sludge-forming
tendencies of lube oils.
Polymerization catalysts
Catalyst for air-blowing of
asphalt
(Continued)
98
-------
TABLE F3-11. Continued
Chemical
Uses
Potassium Hydroxide
Potassium Phosphate
Propane
Sodium Carbonate (Soda Ash)
Sodium Hydroxide
(Caustic Soda)
Sodium Hypochlorite
Sodium Phenolate
Sodium Plumbite
Sulfur Chlorides
Sulfur Dioxide
Sulfuric Acid
Toluene
Trichloroethylene
Remove acids from petroleum
Remove hydrogen sulfide from gas
Solvent extractions-deasphating,
dewaxing, and decarbonizing
Neutralize acids in processing
streams
Remove acidic substances
Sweeten gasoline
Remove hydrogen sulfide from
gasoline
Stabilize color of gasoline
"Doctor sweetening" agent to convert
mercaptans to disulfides
Solvents
Extract aromatic hydrocarbons and
sulfur-bearing compounds from
paraffins and naphthenes
Improve viscosity index and remove
waxes from lube oils
Remove aromatics from kerosene
Remove or dissolve resinous and
asphaltic materials and sulfur
Remove waxes from lube oils
Extract carbon- and sludge-forming
constituents of lube oils and
increase their viscosity index
Sourcer Reference 1
99
-------
Several processing steps result in the production of
refinery gases. The exact composition of a gas stream varies
with its origin. Table F3-12 summarizes the compositions of
the gas streams from the various processes.
One step in the production of high-octane gasoline is
the catalytic reforming of hydrocarbons in the C? to Ci0 range.
The gas produced contains hydrogen and hydrocarbons in the Ci
to Ci, range. The hydrogen is usually recycled and the Cs/Ci,
fraction added to the liquified petroleum gas stream.
Catalytic cracking of heavy oils produces saturated
and unsaturated hydrocarbons. Acetylenes may also be produced
with severe cracking. Naphtha or gas oil may be thermally
cracked to produce ethylene and propylene. Additional gases are
produced by the coking, visbreaking, and hydrocracking of heavy
oils.
3.2.1.2 Liquified Petroleum Gas--
Liquified petroleum gas, also called LP gas, LPG, or
"bottled gas", is a mixture primarily of propane and butane,
both of which are considered potentially hazardous. It may be
sold either as a mixture or as individual gases. It is stored
and transported as a liquid; liquefaction is achieved either by
compression or refrigeration. Only about one-fourth of the LP
gas used in this country is obtained from petroleum; the remain-
der is obtained from natural gas.1
Most LP gas is used as fuel either for heating or for
motor vehicles. However, much of the butane from refineries is
used as petrochemical feedstock.
100
-------
Page Intentionally Blank
-------
TABLE F3-L2. COMPOSITIONS OF TYPICAL REFINERY GASES
Composition
% wt
H,
Ci
C2
C2 =
C3
Cs =
Cli
€„ =
Gas yield,
7. wt (on
feed to unit)
Primary
Fractionator
Gas
--
8.5
15.4
--
30.7
--
45.9
--
9.2
Catalytic
Reformer
(Power former)
Tall Gas
1.5
6.0
17.5
--
31.5
._
43.5
--
20.0
Steam Cracker
Fluid Cat Hydrocracker
Cracker Isoniax Tail Naphtha Gas Oil Hydrofiner
Tall Gas Gas Li&ht Ends Tail Gas
0.6 1.4 1.2 1.3 3.0
7.9 21.8 17.4 19.6 24.0
11.5 4,4 7.0 3.5 70.0
3.6 -- 33.3 38.0
14.0 15.3 0.7 1.0 3.0
16.4 -- 27.0 19.7
21.8
57.1 13.* ie.9
24.2
16.5 14.5 72.8 59.6 3.8
Source: Reference 2
-------
3.2.1.3 Aviation Fuels--
Aviation Gasolines—Aviation gasolines consist almost
entirely of hydrocarbons. In general, these fuels have a high
percentage of isoparaffins and smaller percentages of naphthenes
and aromatics. A typical ultimate analysis is 84 weight percent
hydrogen, 15 percent carbon, and some sulfur, lead, bromine,
nitrogen, and oxygen from dyes and antiknock additives. Most
constituents of aviation fuel are considered potentially
hazardous.2
Knocking in an automobile engine can be a minor
problem; in aircraft engines it can be extremely serious. To
prevent the problem, a small amount of tetraethyl lead is added.
The lead oxide which could be formed from the tetraethyl lead
would form deposits in the combustion zone of the engine and
cause fouling of spark plugs and valves, so ethyl bromide is
added to form volatile lead bromide. The tetraethyl lead and
ethyl bromide must be in an exact 1:1 ratio because excess
bromide would be corrosive to engine parts. An identifying dye
and kerosene are added with the tetraethyl lead/ethylene
bromide mixture. Fuels with octane ratings well above 100 are
possible.2
Otiher properties of the fuel mixture must be geared
to specific requirements. For instance, if the fuel is not
volatile enough, it will be difficult to start a cold engine;
if it is too volatile, it will vaporize in the tanks, particu-
larly at high altitudes. A boiling range of from about 85°F
to 300°F and a Reid vapor pressure of between 5.5 and 7.0 psi
(vapor pressure at 100°F) is desirable. The freezing tempera-
ture must be -76°F or below. Average viscosity is 0.75 centi-
stokes.at 32°F.2
102
-------
. ;-_- A large proportion of the lower paraffins and isoparaf-
fins is desirable to minimize carbon deposits and to maximize
heat energy. Antioxidants are added to retard the formation of
gum and the precipitation of lead compounds during storage.
Jet Fuels—An approximate distribution of the hydro-
carbons in jet fuels is given in Table F3-13. Aromatics are
carefully limited, sometimes to even lower values than given,
because they act as a solvent on fuel lines. Polynuclear
aromatics are sometimes eliminated. Olefin and sulfur contents
are limited, but not nitrogen and oxygen.
TABLE F3-13. APPROXIMATE DISTRIBUTION OF HYDROCARBONS
IN JET FUELS
Paraffins: 33-6170 volume percent
Olefins: 0.5-5% volume percent
Naphthenes : 10-4570 volume percent
Total Aromatics: 12-25% volume percent
Source: Reference 2
Additives are added to the fuel to control oxidation,
to chelate any copper that may be present from refining, to
ensure that any water dissolved in the fuel will not freeze, to
increase conductivity and thus reduce static electricity, and
to inhibit corrosion. Most constituents of jet fuel are con-
sidered potentially hazardous.
. The boiling range of jet fuel is about 300°F to 460°F;
maximum viscosity is usually about 5 centistokes at -30°F.
103
-------
Freez-ing temperature must be -76°F or below, though this limit
may be raised as high as -40°F in some cases.
3.2.1.4 Automobile Gasoline--
Gasoline is defined as a petroleum fuel for use in
reciprocating, spark-ignition, internal combustion engines. It
is a complex mixture of hydrocarbons, mostly in the d, to Ci2
range, which distill between 85°F and 410°F. Gasolines from
different refineries may vary widely in exact composition
according to the processes used at the refinery. A summary of
the main components of gasoline and their sources is given in
Table F3-14.
In general, gasoline must meet the following criteria:
A. Burn smoothly without knocking,
B. Evaporate readily enough that a combustible
fuel-air mixture is supplied when the engine is
started cold and that a considerable portion
is vaporized in the intake manifold when the
engine is run warm,
C. Evaporate slowly enough that it will not boil
in the fuel pump or fuel lines, and
D. Evaporate completely with no residue.
To meet these criteria, each refinery has its own blending
scheme. Additives that may be added include antiknock com-
pounds , anti-icing additives, anti-oxidants, metal deactivators,
carburetor detergents, anti-corrosion additives, and others.
These are discussed in Section 3.I.2.2.1'*
104
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TABLE F3-L4. MAIN COMPONENTS OF GASOLINE
Components
Source
Boiling
Ranp.e, °F
Remarks
Paraffinic
Butane
leopentane
Alkylate
Isomerate
Straight-run
Naphtha
Hydrocrackate
Olefinic
Thermal Reformate
Catalytic Naphtha
Steam Cracked
Naphtha
Polymer
Aromatic
Catalytic
Reformate
Crude oil distillation 30
Conversion processes
Crude oil distillation 81
Conversion processes
Isomerization of
n-pentane
Alkylation process 100-300
Isomerization process 100-160
Crude oil 90-200
distillation
Hydrocracking 100-390
process
Thermal reforming 100-390
Catalytic cracking 100-390
Steam cracking 100-390
Polymerization of 140-390
olefins
Catalytic reforming 100-390
Widely used in proportions
up to 10%.
Widely used as high-octane,
high-volatility component.
Used widely in aviation gaso-
line, but less frequently in
motor gasoline.
Relatively little used at
present. Excellent anti-
knock properties under severe
engine conditions.
Widely used low-octane compo-
nent .
Heavy products used as feed
for catalytic reforming. Con-
tains also aromatics.
Obsolescent process.
Widely used component, par-
ticularly In premium gasoline.
By-product of chemical processes.
High-octane component but not
widely used.
Most widely used high-octane
component of gasolines.
Source: Reference 2
-------
3.2.1.5c- Diesel Fuel--
Diesel fuel for commercial use is divided into three
categories according to ASTM Specifications:
• Grade 1-D: volatile fuel oils from kerosene
in the intermediate distillates.
• Grade 2-D: distillate gas oils of lower
volatility.
• Grade 4-D: more viscous distillates and blends
of these distillates with residual
fuel oils.
The exact composition of a diesel fuel varies with
the refinery processes used and with product requirements.
Most constituents are considered potentially hazardous.1
3.2.1.6 Heating Oils--
Heating oils include kerosene, distillate oils, and
residual oils. The only uniform requirement appears to be that
they be free of acid, grit, and fibrous or other foreign matter
likely to clog or injure burner or valves. Other more specific
requirements may vary.
Kerosene is a broad term for the virgin refinery
product which has a boiling range of 350°F to 550°F and an API
gravity of 43° to 45°. Characteristics of three grades of
distillate oil are given in Table F3-15 and of two grades of
residual oil in Table F3-16. Most constituents of heating oils
are considered potentially hazardous.1
106
-------
TABLE F3-15. CHARACTERISTICS OF THREE GRADES OF DISTILLATE
- HEATING OILS
Property Grade 1 Grade 2 Grade 4
Gravity, "API
Viscosity at 100°F, cs
Sulfur, wt %
Rair.sbottom carbon residue,
wt %
Distillation, °F:
Initial boiling point
107C point
50% point
Final boiling point
42.
1.
0.
0.
349
390
437
533
6
79
071
052
34.
2.
0.
0.
370
432
499
629
9
61
249
116
21
15
0
3
422
496
674
754
.2
.41
.77
.30
Source: Reference 1
TABLE F3-16. CHARACTERISTICS OF TWO RESIDUAL-TYPE
HEATING OILS
Property
Gravity, °API
Viscosity :
Kinematic at 100°F, cs
Furol at 122°F, sec
Sulfur content, %
Ramsbottom carbon residue on
100% sample, °L
Ash, %
Water and sediment, vol 70
Grade 5
17
60
25
1
6
0
0
.1
.2
.8
.07
.7
.035
.16
Grade 6
12.
170.
1.
10.
0.
0.
3
2
33
7
41
15
Source: Reference 1
107
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3.2.1.7— Industrial Fuels--
The term "industrial fuels" as used here refers to gas
oils and fuel oils. These may be straight-run, cracked, or a
blend of both.2
Gas Oil--Gas oil is an intermediate between kerosene
and light lubricating oil. It has no residual components and is
low in sulfur. It is commonly used as a thinning agent for
reducing the viscosity of heavy fuel oil without lowering the
flash point too much. The characteristics of individual gas
oils depend on the nature of the crude and its subsequent pro-
cessing.
Fuel Oil—Fuel oils include heavy residual oils and
such oils blended with varying amounts of gas oil to improve
viscosity. Specific gravity may range from 0.92 to l.O.2
Fuel oils for specific markets are blends of several
components. Table F3-17 shows typical analyses of three example
fuel oils. Most constitutents are considered potentially
hazardous.
3.2.1.8 Solvents, Thinners and Diluents (Industrial
Naphthas)--
Almost all industrial naphthas are clear, water-white
liquids which are immiscible with water but completely miscible
with other organic substances. They may be pure hydrocarbons
such as benzene, toluene, xylene, ethylbenzene, hexane, or
cyclohexane, or they may be blends of varying proportions of
paraffins, cycloparaffins and aromatics. Solvents usually con-
tain no olefins.
108
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TABLE F3-17. TYPICAL ANALYSES OF INDUSTRIAL FUEL OILS
Light
Specific gravity at 15/59°C
Kinematic viscosity at
48.4°C, cSt, max -
Calorific value, gross,
Btu/lb 18
Ultimate analysis, % wt
C
H
S
Ash
N,0, etc (by difference)
C:H ratio
Flashpoint, closed,
Pensky -Mar tens , °C, min
0
12
,700
84
12
2
0
0
7
66
.922
.5
.9
.0
.57
.06
.47
.1
Medium
0
30
18,500
84
11
3
0
0
7
66
.948
.9
.4
.19
.11
.40
.4
Heavy
0.
70
18,190
84.
11.
3.
0.
0.
7.
66
977
5
1
84
12
44
6
Source: Reference 2
109
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. i- There are some standard solvents with properties
specified by the ASTM Committee on Petroleum Products and
Lubricants. Properties of three such solvents are given in
Table F3-18. Most common constituents of industrial naphthas
are considered potentially hazardous. Benzene in particular
has been cited for its extremely toxic properties.
3.2.1.9 Lubricants--
An API survey, a few years ago, reported over 1100
lubricating oils and nearly 300 grease products made by U.S.
companies. The great variety of applications makes a standard
set of specifications impossible.2
Lubricants are of two types: automotive and indus-
trial. Automotive lubricants include crankcase oils, transmis-
sion and axle lubricants, and fluids for hydraulic torque
converters and fluid couplings such as automatic transmissions.
Industrial lubricants include those required for the power plant
and those for production equipment.
Viscosity is generally the single most important
property of lubricants. Other important properties are API
gravity, carbon residue, cloud and pour points, color, corrosive
tendencies, foaming tendencies, sulfur content, saponification,
and neutralization number.
Lubricating oils are obtained from the vacuum distil-
lation of the residue from atmospheric distillation. Not all
crudes produce acceptable lube oils; therefore several refining
steps are usually necessary.
110
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TABLE F3-18. PROPERTIES OF THREE INDUSTRIAL SOLVENTS
- -I-" FOR WHICH SPECIFICATIONS HAVE BEEN
PREPARED BY ASTM COMMITTEE D-2 ON
PETROLEUM PRODUCTS AND LUBRICANTS
PETROLEUM SPIRITS (MINERAL SPIRITS) . D235
Appearance: Clear and free of suspended matter
and water
Color: Water white, not darker than No.
21 Saybolt Chromometer
Flash point, nin: 100°F
Blackening: Shall not blacken or corrode clear.
metallic copper in 30 cin. at boil-
ing point of spirits.
Distillation:
7. recovered at 350°?,
min 50
End point 410°F
Acidity of distillation
residue Neutral
HEAVY PETROLEUM SPIRITS (HEAVY MINERAL SPIRITS), D965*
Appearance: Clear and free of suspended natter
and water
Color: Water white, noc darker than No.
21 Saybolt Chronometer
Flash point, min: 125°F
Blackening: Shall not blacken or corrode clean
metallic copper in 30 min at boil-
ing point of spirits
(Continued)
111
-------
TABLE F3-18. Continued
Discillation:
Initial boiling point,
min 340°F
95% (by volume) 460°F
End point, taax 485°F
Acidity of distillation
residue: Neutral
*Tentative.
STODDARD SOLVENT, D 48^
Color:
Corrosion at 212'F
Doctor test:
Sulfuric-Acid
absorption, max, *
Flash point, min. °F:
Distillation:
Percentage recovered
at 350*F, min
Percentage recovered
ac 375"F, min
End point, max. *F
Residue, max, 1,
Acidity:
Hater white or not darker than 21
Not more than extreaely slight cis-
coloraticn of the copper test strip,
or shows no greater corrosion than
a mutually approved reference scrip
Negative
100
50
90
410
1.5
No acid reaction to methyl
orange shown by residue from
distillation
So-jrce: Reference 1
112
-------
. i- The major hydrocarbons in lubricating oils are
paraffinic compounds, polycyclic and fused ring hydrocarbons
based on naphthenes, and mono- and polynuclear aromatics.
Additives are added to improve the properties of the oil.
Lubricating greases are essentially lubricating oils
with thickening agents added. The oil used usually has a high
viscosity index. The major thickening agents are soaps of
aluminum, barium, calcium, lithium, sodium, and strontium.
Solid fillers such as asbestos, graphite, metal oxides, metal
powders or flakes, or metal sulfide may be included and additives
similar to those used for lubricating oils are also used.
3.2.1.10 Waxes--
Waxes are separated from crude fractions often more
as a measure to improve the quality of the fraction rather than
to recover the wax. The properties of the waxes vary with the
fraction they are recovered from. A comparison of the major
types of wax is given in Table F3-19.
Paraffin waxes are by far the most common waxes used
today. They are composed almost entirely of normal alkanes with
some isoalkanes, cycloalkanes and only a trace of aromatics.
Refined paraffin wax has a light color and contains less than
0.5 percent oil.l'2
Less is known about the microcrystallic waxes. They
have molecular weights of about 600-800 and appear to contain
branched alkanes, aromates, polycycloparaffins, and monocyclo-
pentyl, monocyclohexyl, and dicyclohexyl paraffins in addition
to the normal alkanes.1'2
113
-------
TABLE F3-19 COMPARISON OF MAJOR WAX TYPES PRODUCED
IN THE UNITED STATES
Wax
No. of Melting Viscosity
Charac- Carbon Point, at 210°F,
teristic Atoms °F SSU
Source: Reference 1
Crystals
Paraffin Brittle 18-56 122-140 40
Micro-
crystalline
Motor Oil Brittle 26-42 145-170 50
Residual Flexible 36-70 145-175 65-100
.Tank bottom Hard 40-70 180-200
Plates
Needles
Small needles
Very small
needles
-------
. i.- Petrolatum is a mixture of macrocrystalline waxes and
oil. It is described as a colloidal system in which the solid
hydrocarbons are the external phase and the liquid components
are the internal phase. Petrolatum is a semisolid at room
temperature; its melting point ranges from 110 to 175°F.'
Petroleum waxes are considered to be nonhazardous if
they are pure.
3.2.1.11 Asphalt--
Asphalt cement is the material remaining after the
removal of light and heavy distillates from asphaltic crudes.
It is usually mixed with distillates in varying proportions to
obtain materials for specific purposes.
Cutback asphalts contain lighter distillates such as
naphthas, gasoline or kerosene. They may be medium- or rapid-
curing.
Emulsified asphalts are emulsions of asphalt cement
with chemically treated water. Road oils contain a large pro-
portion of oil which is either allowed to remain during dis-
tillation or is added during processing.
"Blown asphalt" may be produced by blowing air through
a residual oil at temperatures usually about 400 to 600°F.'
Asphalts of varying characteristics may be produced. Air may
also be blown through asphalt to remove hydrogen as water.
Partially dehydrogenated molecules are polymerized.
115
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3.2.1.1-2 Petrochemicals- -
A widely accepted definition of a petrochemical reads
"a chemical compound or element recovered from petroleum or
natural gas, or derived in whole or in part from petroleum or
natural gas hydrocarbons and intended for chemical markets."
There are hundreds of petrochemicals manufactured today.
Petrochemicals are grouped according to their chemical
composition and structure as aliphatic, aromatic or inorganic.
Most aliphatic petrochemicals are made from Ci to Ci, hydrocar-
bons. Important aromatic petrochemicals are benzene, toluene,
xylene, phenol and styrene. Typical inorganic petrochemicals
are sulfur, ammonia and its derivatives, and carbon black.1
More than half of the petrochemicals are aliphatic.
A list of 95 commercial petrochemicals, their important proper-
ties and the raw material they are derived from is given in
Petroleum Processing Handbook edited by Wm. F. Bland and Robert
L. Davidson, McGraw-Hill, New York, 1967.
3.2.1.13 Carbon Black and Petroleum Coke--
Carbon black, though a petrochemical, is often con-
sidered separately. It is the blackest pigment available to
industry.
Carbon black is a finely divided semigraphitic form
of carbon. It can be produced by the partial combustion of oil
in a furnace. Its properties vary with particle size and the
related surface area.
116
-------
. i; Coke is a by-product of the thermal cracking of
reduced crudes and residuums. Its composition varies with the
feedstock. Sulfur and metal content are important considera-
tions.
To make coke more suitable for certain purposes, it is
calcined. Calcination is a high-temperature treatment in which
the carbon-hydrogen ratio is increased from about 20 to more
than 1000. Volatile material is also driven off and the coke is
dehydrogenated.J
3.2.2 Intermediate Products
An intermediate product is considered here to be any
material which is produced by one refinery process and subse-
quently used in another. Most intermediate products used as
feedstocks for major processes have been discussed previously in
this section either as components of the raw crude or as final
products. Others will be discussed here. Table F3-20 is a list
of major processes and their major feedstocks.
Light ends include the light gases (methane, ethane,
and propane) from the crude plus the light gases from a number
of processes. Some light ends come from cracking operations;
therefore, olefins and isoparaffins in the C2 to d, range may
also be present in a light ends stream. The light ends stream,
like many other intermediate streams, may be sweet or sour,
depending on whether or not sulfur compounds have been removed
or made less objectionable. The olefin stream is simply the
C2 to C4 olefins which have been separated from the other light
ends.
117
-------
TABLED. F3-20.
MAJOR PROCESSES AND THEIR MAJOR
FEEDSTOCKS
Process
Major Feedstocks
Atmospheric Distillation
Vacuum Distillation
Residual Oil Hydro-
desulfurization
Asphalt Blowing
Lube Oil Processing
(Dewaxing)
Coking
Gas Oil Hydrodesulfurization
Catalytic Hydrocracking
Fluid Catalytic Cracking
Kerosene Hydrodesulfurization
Naphtha Hydrodesulfurization
Catalytic Reforming
Isomerization
Gas Processing
Alkylation
Sulfur Recovery
Crude oil
Topped Crude
Residual Oil
Raw Asphalt
Hydrotreated Lube Oil
Residual Oil
Vacuum Oil
Gas Oil
Gas Oil
Kerosene
Vacuum Oil
Deasphalted Oil
Kerosene
Gas Oil
Kerosene
Sour Naphtha
Gasoline
Sweet Naphtha
Sweet Light Ends
Sweet Naphtha
n-butane
Sweet Light Ends
Isobutane
Olefin Gases
Sour Light Ends
118
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3.3 ,. REFERENCES FOR SECTION 3
1. Bland, William F., and Robert L. Davidson, eds. Petroleum
Processing Handbook. McGraw-Hill Book Company, New York,
1967.
2. Hobson, G. D., ed. 'Modern Petroleum Technology, 4th Ed.
Applied Science Publishers, Ltd., Essex, England, 1973.
3. Qualitative and Quantitative Aspects of Crude Oil Composi-
tion. U.S. Bureau of Mines Publication No. 642. 1968.
4. Powell, D., et al. Development of Petroleum Refinery Plot
Plans. EPA-450/3-78-025. Pacific Environmental Services,
June 1978.
5. Bombaugh, K. J.( et al. Sampling and Analytical Strategies
for Compounds in Petroleum Refinery Streams. Final report,
2 volumes, PB 251 744 (vol. 1); PB 251 745 (vol. 2); EPA-
600/2-76-012, Radian Corp., Austin, TX, January 1976.
6. Bell, H. S. American Petroleum Refining, 4th Ed. D. Van
Nostrand, Princeton, 1959.
7. Hunt, R. H., and M. J. O'Neal, Jr. The Composition of
Petroleum. Adv. Pet. Chem. & Refin. 10, pp. 3-34, 1965.
8. Gary, James H., and Glenn E. Handwerk. Petroleum Refining.
Marcel Dekker, Inc., New York, 1975.
9. Rossini, F. D., and S. S. Shaffer. API Research Project 6--
Analysis, Purification, and Properties of Petroleum Hydro-
carbons. API Proc. 34 (Sect. 6), 14, 1954.
10. Tye, Russel, et al. Carcinogens in a Cracked Petroleum
Residuum. Arch. Env. Hlth., 13:302, 1966.
11. Carruthers, W. The Constituents of High-Boiling Petroleum
Distillates, Pt. 3. Anthracene Homologues in a Kuwait Oil.
J. Chem. Soc., 1956.
12. McKay, J. F., and D. R. Latham. Polyaromatic Hydrocarbons
in High-Boiling Petroleum Distillates. Isolation by Gel
Permeation Chromatography and Identification by Fluores-
cence Spectrometry. Anal. Chem. 45(7):1050-55 , 1973.
119
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13. Hirisch, D. E., et al. Compound-Type Separ-a-tion and
Characterization Studies for a 370° to 535°C Distillate of
Wilmington, CA Crude Oil. Bartlesville Energy Research
Center, Bartlesville, Oklahoma, 1974.
14. Tyler, A. L. , and J. A. Apps'. A Simple Model for Predict-
ing Transient Responses in Dump Leaching Operations.
Presented at the 78th National AIChE Mtg., Salt Lake City,
Utah. University of Utah, 1974.
13. Thompson, C. J., et al. Analyzing Heavy Ends of Crude.
Hydrocarbon Proc., 52(9):123, September 1973.
16. Dooley, J. E. , et al. Analyzing Heavy Ends of Crude.
Hydrocarbon Proc., 53(4):93, April 1974.
17. Yen, T. F. The Role of Trace Metals in Petroleum. Ann
Arbor Science Publishers, Inc., Ann Arbon, MI. 1975.
120
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SECTION 4
REFINERY PROCESS TECHNOLOGY
Section 4 contains a characterization of various major
refining processes or operations. All of the processes which
have been utilized in the development of a "hypothetical refinery"
(discussed in Section 2 of this appendix) are discussed in this
section. In addition, several other common processes not
included in the hypothetical refinery are also described here.
The discussion of each of the processes described in
this section includes the following types of information:
• A description of each process,
• A description of process technology, and
• A discussion with appropriate estimates for
atmospheric emissions from each type of
process.
A process description is provided for each unit
discussed in Section 4. Included is information on the function
of the process, a description of significant differences between
processes of the same general type (e.g., sulfuric versus hydro-
fluoric acid alkylation), and a discussion on how the process
is integrated into the overall refining scheme.
121
-------
A detailed description of the operation of each process
is contained in the subsection titled "process technology".
Here, specific details concerning process operations, major
processing variables, feed and catalyst characteristics, recent
process innovations, etc., are described.
A subsection on atmospheric emissions has been provided
for most of the processes discussed in Section 4. Emissions
sources can include process emissions such as those from fired
heaters and catalyst regeneration operations, and fugitive
emissions from various process equipment.
Three types of information necessary to estimate total
atmospheric emissions from refinery process units have been
studied over the course of this program:
• Fugitive hydrocarbon emission factors for each
type of source have been determined as a
function of the type of process fluid handled
by the source.
• An estimate of the number of fugitive sources
within each type of process unit was obtained,
based on field source counts.
• The distribution of these sources between
the process fluid groups used in the calcu-
lation of emission factors was investigated.
122
-------
The development of emission factors for fugitive
sources is a major result of this program and the methods used
to establish these emission factors are well documented in
numerous places in this report. Conversely, the estimates for
the number and process fluid distribution of fugitive sources
in specific process units are not as reliable.
Where possible, two estimates of the number of fugitive
sources are provided for each type of process unit. The first
estimates represent actual field source counts, or estimates
based on these counts, which were developed as a part of this
program. The second set of estimates were developed in a
published EPA report, The Development of Petroleum Refinery Plot
Plans (Reference 8).
The estimates developed during this program were based
on field source counting in 19 process units. This is an
insufficient amount of data to establish accurate average source
counts for individual process units. However, it was determined
that the number of sources within similar process units can
vary considerably from one refinery to the next. Further, these
variations in the number of sources do not correlate with the
capacity of the counted process units. The methods used to
estimate the number of fugitive sources and the distribution
of these sources between the four process fluid groups are
described in detail in Section 4.1.1.3.
The source count estimates provided in Reference 8 were
based on counts developed from process flow diagrams. In particu-
lar, the numbers of pumps, compressors, and control valves were
counted from flow diagrams or estimated, based on results from
process units containing similar vessel types. From these
counts T- ratios were used to estimate the number of block valves,
123
-------
flanges:, and screwed fittings. Relief valve and sample valve
estimates were obtained by other means. Details of the methods
used to estimate the sources counts presented in Reference 8
are also included in Section 4.1.1.3.
Both sets of estimates described above have been used,
where possible, to allow the calculation of a range of emissions
from each of the process units. It is evident that, in a number
of cases, the two sets of estimates are quite different. This
is probably a result of real differences between the number of
sources in actual process units as well as differences resulting
from the methods used to obtain these estimates. It is important
that both sets of source counts be considered, at best, rough
estimates. Estimates for total fugitive emissions from actual
process units should be based on the number and distribution
of sources actually present in that particular process unit.
4.1 SEPARATION PROCESSES
Crude oil separation describes those processes which
separate crude oil into a variety of intermediate products.
These products become feedstock for additional downstream refin-
ery processing units. Distillation streams are typically
classified according to boiling range, as shown in Table F4-1.
Higher efficiencies and lower costs are achieved if
crude oil separation is accomplished in two steps: 1) fraction-
ating the total crude stream at atmospheric pressure; then 2)
feeding the high boiling bottoms fraction (topped crude) from
the atmospheric still to a second fractionator operating under a
high vacuum. Sections 4.1.1 and 4.1.2 describe process operation
and process emissions for atmospheric and vacuum distillation
units. ^
124
-------
TABLE "F4-1. TYPICAL BOILING RANGES FOR CRUDE "OIL FRACTIONS
Fraction Boiling Range, °F
Butanes and lighter gases 80
Light straight run gasoline (LSR) 80-220
Naphtha (heavy straight run gasoline) 180-400
Kerosene 330-540
Light gas oil 420-640
Atmospheric gas oil 550-830
Vacuum gas oil 750-1050
Vacuum reduced crude 1000
The aromatic hydrocarbon content of reformate streams
is quite high. Aromatics extraction units are used to separate
and concentrate these aromatic compounds for use in gasoline
blending or petrochemical operations. Aromatics extraction is
discussed in Section 4.1.3.
4.1.1 Atmospheric Distillation
Nearly all incoming crude oil feed must pass through a
refinery's atmospheric distillation unit before it can be further
processed. The purpose of atmospheric distillation is to sepa-
rate the crude's hydrocarbon components into fractions by means
of distillation and steam stripping.
Because it performs a function basic to refining,
almost all refineries will have at least one atmospheric distil-
lation unit. United States atmospheric distillation capacity
was 17._2 x 106 bbl/day in 1978. Approximately 14.8 x 106 bbl/
day of crude was processed in 1978.
125
-------
4.1.1.1 Process Description--
The atmospheric distillation process is outlined in
Figure F4-1. After desalting, the crude oil is pumped through
a series of heat exchangers in which its temperature is elevated
by heat recovered from product and reflux streams. The preheated
crude is then charged to a direct-fired furnace where additional
heat is supplied to achieve partial vaporization of the crude
petroleum. Both the liquid and vaporized portions are charged
to the atmospheric fractionator at a temperature of about
650-700°F. At temperatures above this range, thermal decomposi-
tion ("cracking") of the crude may occur. The carbon produced
by cracking fouls equipment and has a deleterious effect on
properties of some distillation products. In addition, some of
the more valuable liquid products are cracked at high tempera-
tures to form less valuable gases.
The crude charge is separated into several petroleum
fractions within the atmospheric fractionator. A stream of
naphtha and lighter material is condensed and taken from the
tower overhead. This stream is essentially light straight run
(LSR) gasoline containing some propane, butanes, and almost all
the pentane-and-heavier components in the tower overhead vapor.
Some of the condensate is returned to the top of the tower as
reflux, and the remainder is sent to other refinery processes
or product storage.
Several liquid side stream fractions are withdrawn
from the fractionator at different elevations within the tower.
These streams are charged to the side stream product strippers,
where lighter hydrocarbons are stripped and returned to the
fractionation tower. The stripping medium is typically steam,
light petroleum gases, or reboiler vapors. In addition to the
126
-------
CRUDE CHARGE
SALT WATER
GAS TO LPG
RECOVERY
WATER
LSR GASOLINE
TO TREATING
NAPHTHA
GAS OIL
TOPPED CRUDE TO
VACUUM TOWER
70-1511-1
Figure F4-1. Atmospheric distillation.
-------
side stream strippers, the atmospheric fractionator has a bottoms
stripping zone in which lighter hydrocarbons are stripped from
the residual product.
The stripping medium and stripped hydrocarbons are
vented back to the vapor zone of the fractionator at a point
above the corresponding side-stream withdrawal point.
The fractions withdrawn from the atmospheric
fractionator are progressively heavier as they are removed at
successively lower points. The vaporization end point of the
heaviest sidestream product closely corresponds to the crude's
temperature when charged to the fractionator.
Fractionator bottoms (topped crude) is the heaviest
petroleum fraction in the crude. Because this fraction cannot
be further separated at atmospheric pressure, it is charged to
the vacuum distillation unit.
Relative product amounts from crude atmospheric distil-
lation vary with the type of crude being processed. Variances
in product amount range from 10 to 50 percent.
In general, the intermediate products from atmospheric
distillation include liquid petroleum gases (LPG), light straight
run gasoline (LSR), naphtha, kerosene, distillate or diesel oil,
gas oil, and topped crude. The naphtha may be blended into
motor fuels or any of several other refinery products, or further
processed to improve octane rating and/or reduce sulfur content.
The kerosene may be chemically sweetened or hydrogen treated and
sold directly or blended with other petroleum products. The
distillate or diesel oil may be sold directly as diesel fuel or
fuel oil. It may also be hydrogen treated, hydrocracked, cata-
lytically cracked, or blended with other refinery products. Gas
128
-------
oil is -not usually used directly, but may be processed similarly
to distillate oil. Although the topped crude fraction is usually
the feed for the vacuum distillation process, it may also be sold
for fuel, blended into fuels, hydrogen treated, catalytically
cracked, or sent to a coking unit.
4.1.1.2 Process Technology--
Process Conditions, Fuel, and Utility Requirements--
Typical operating parameters for an atmospheric distillation unit
with a capacity of 24,000 bbl/day are listed below:
• Pressure: Atmospheric
• Temperature: 250°F - at top of fractionator
700°F - at bottom of fractionator
• Electricity: 4.1 kW/bbl
• Thermal Energy: 105 Btu/bbl
• Steam: 50 Ib/bbl
• Process Water: 50 gal/bbl
Process Equipmenc--The major pieces of process equip-
ment used in atmospheric distillation include heat exchangers,
a direct fired furnace, an atmospheric fractionation column,
and side stream product strippers. A brief description of each
of these items follows.
129
-------
- ~~ Heat exchangers are employed for a number of uses:
feedstock preheating and vaporization, condensation of overhead
vapor, and cooling liquid products before discharge to further
processing or storage. The most common type of heat exchanger
used in refineries is the shell-and-tube type, which permits
high heat transfer rates by allowing high fluid velocities and
countercurrent flow between tube side and shell side fluids.
Commonly used shell-and-tube exchangers include fixed tube sheet,
floating head, and hairpin-type tube units.
Heat exchangers may be air or water cooled. Initial
capital costs and space requirements for air-cooled units are
usually higher than those for water-cooled units. However,
operating costs for water-cooled units are often higher because
of higher costs for cooling water, wastewater disposal, and tube
maintenance.
A direct-fired furnace ("pipestill") is used to heat
the crude charge so that it is partially vaporized when it enters
the fractionation tower. A typical furnace consists of an array
of tubes arranged in a refractory lined shell that may be
cylindrically- or box-shaped. Burners located in the heater
shell wall provide heat, which is transferred by both radiation
and convection to the crude oil inside the tubes.
The type of heater used depends on the heat load as
well as the thermal efficiency and pressure drop of the fluid to
be heated. The high heat loads required in crude distillation
units necessitate the use of box furnaces, which are usually
designed for heat absorption rates greater than 150 x 106 Btu/hr.
130
-------
^.- The atmospheric fractionator is a vertical cylindrical
column 'in which vapor is bubbled through the liquid retained on
a vertical array of horizontal trays. The fractionation tower
normally contains 30 to 50 fractionation trays. Five to eight
trays are needed for each sidestream product, and the same number
is required above and below the feed point.
Several different types of trays are used in atmos-
pheric distillation towers: bubble cap trays, perforated trays,
grid trays, and valve trays. Valve and perforated trays are most
widely used in atmospheric distillation units. Bubble cap trays
are highly flexible and efficient but have a high initial cost.
Perforated trays are lower in cost, but are also lower in
efficiency at low vapor loadings. Grid trays, while economical,
are relatively inflexible in handling varying loads. Valve trays
are considerably less expensive than bubble cap trays, and are
also capable of operating at high efficiencies over a wide range
of loadings.
Any number of sidestream stripping towers may be
employed in an atmospheric distillation unit, depending on the
number of distillation cuts being taken. Each stripping tower
contains four to ten trays. Steam is introduced under the bottom
tray.
4.1.1.3 Atmospheric Emissions--
Emissions sources from atmospheric distillation
operations include:
• Process heater flue gas emissions, and
• Fugitive emissions.
131
-------
— Process Heater Emissions--Emission factors for various
polluta'nts from fired heaters are given in Table F4-2. A
typical atmospheric distillation unit will require a process
heater to raise the temperature of the crude oil prior to its
entry into the fractionator. Additional process heat is re-
quired to reboil the fractionator bottoms.
Total emissions from an atmospheric distillation unit
process heater are also given in Table F4-2. These figures are
given as pounds of pollutant per thousand barrels of feed using
a typical heat input requirement of 1.0 x 10s Btu per barrel of
feed.
Fugitive Emissions—Fugitive emissions can occur from
a variety of sources including valves, pumps, compressors,
flanges and other fittings, relief valves, and drains. Hydro-
carbon emission rates for each individual source are quite low.
However, total fugitive emissions are significant because of the
large number of such sources within the refinery.
Total estimated fugitive non-methane hydrocarbon emis-
sions from a typical crude distillation unit are given in Table
F4-3. The listed emission factors were determined as a part of
this program. Additional information on these emission factors,
including a complete discussion on their derivation and the con-
fidence intervals for each source category, are contained in
Appendices B and C of this report.
Estimates for the number of sources within each source
category were developed from:
• Actual field source counts conducted during the
course of this program, and
132
-------
- 1ABLE F4-2.
TYPICAL EMISSIONS FROM ATMOSPHERIC
DISTILLATION UNIT PROCESS HEATERS
EPA Emission Factor
(lb/103 gal-oil fired)
(lb/106 scf-gas fired)
Total Emissions
(lb/103 bbl of
crude oil feed)
Oil Fired Beaters
Participates
- Distillate oil
- Residual oil
Grade 4
Grade 5
Grade 6
Sulfur Dioxide0
- Distillate oil
- Residual oil
Sulfur Trioxidec
Carbon Monoxide
Hydrocarbons (as CUt,)
Nitrogen Oxides
(as N02)
- Distillate oil
- Residual oil6
7
10
10(S)+3
142(S)
157(S)
2(S)
5
1
22
22+400 (Uy
1.4
5.0
7.1
101 (S)
112 (S)
3.6
0.71
16
16+286 00"
Gas Fired Heaters
Participates
Sulfur Oxides (as 502)
Carbon Monoxide
Hydrocarbons (as CHi*)
Nitrogen Oxides (as N02)
5-15
0.6
17
3
120-230
0.48-1.43
0.057
1.6
0.29
11.4-21.9
Source: Reference 9
Based on a heat input of l.OxlO5 Btu/bbl of fresh feed with the following
fuel heating values: Oil - 140,000 Btu/gal; Gas - 1050 Btu/scf.
CS = Wt 2 sulfur in the fuel oil
Improper combustion nay cause a significant increase in emissions
^se this emission factor for residual oils with less than 0.52 QJ<.5) nitro-
gen content. For oil with higher nitrogen content (N>0,5}, use emission
factor of 120 lb/103 gal
Based on sulfur content of 2000 gr/10s scf
133
-------
TABLE F4-3.
00
ESTIMATED FUGITIVE NON-METHANE HYDROCARBON EMISSIONS
FROM A TYPICAL CRUDE DISTILLATION UNIT
1
Emissions
Source Type
Valves
(VP
(VP
Open-End
(Sample)
Valves
Pumps (Pump
Seals) (VP
(VP
Drains
Flanges &
Fittings
Relief Valves
Compressors
(Compressor
Seals)
Process
Stream Service (
Classification
Gae /Vapor
Light Liquid
> 0.1 pala 9 100'F)
Heavy Liquid
< 0.1 psla 9 100'F)
Hydrogen Service
Total
All
Light Liquid
> 0.1 psla 9 100'F)
Heavy Liquid
< 0.1 psla 9 100'F)
Total
All
All
All
Hydrocarbon
Hydrogen
Total
Number of Sources
In Process Unit
founts or Estimates Counts or Estimates Emission
From Radian Study From PF.S Study >e Factor, lb/l
89
281
523
0
~59la
-
11(15)
20(28)
31(43)a
69a
£1
3997
6C
1(2)
0
l(2)a
263 - 270 0.059
1663 - 1727 0.024
704 - 703 0.0005
0 0.018
2630 - 2700C
56 - 57b 0.005
26(36)-27(38) 0.25
11(15)-11(15) 0.046
37(52)-38(53)b
0.070
8695 - 8930C 0.00056
0.19
0 1.4
0 0.11
0
Estimated Total
Emissions, ',''
ir Ib/hr
5.25
6.74
0.262
12.3
0.280
3.75
0.690
4.44
2.24
0.0
0.0
25.2
- 15.9
- 41.4
- 0.352
0.0
- 57.7
- 0.285
- 9.50
- 1.29
- 10.8
4.83
- 5.00
1.14
- 2.80
0.0
- 2.80
- 82.6
Physically Counted
Counted From Flow Diagrams
n
Estimated
Reference 8
This FES estimate includes vacuum distillation as part of the crude
distillation unit. Radian estimates for emissions from vacuum
distillation are listed in Section 4.1.2.3 and may be added to the
Radian estimates for atmospheric distillation for comparison to the
PES estimates.
-------
- i • Counts contained in The Development of Petroleum
Refinery Plot Plans, EPA publication number
450/3-78-025.8
The methods used to develop estimates for the number
and process fluid distribution of each source type from data taken
during this program are discussed below. These methods have been
used throughout this appendix wherever estimates of process unit
source counts are provided. The source count data on which these
estimates are based is provided in Section 2.7 of Appendix B
(Detailed results).
Pumps - The number of pumps within a given process unit
was determined from field data collected during this
program. For process units which were not physically
counted, the number of pumps was estimated, based on
a comparison to similar process units or information
from available process flowsheets.
The distribution of these pumps between light and
heavy streams was based on field data taken during
this program or estimated, based on results in similar
process units.
The number of seals (on which total emissions were
based) was assumed to average 1.4 seals per pump. This
value was obtained from the Los Angeles Joint Study10
and is indicated in parenthesis behind the actual
number of pumps in Table F4-3.
Valves - From an analysis of field data, the number of
valves in a process unit was determined to average 41
valves per pump. The Los Angeles Joint Study10 con-
cluded that 23.6 percent of all valves were in gas/vapor
135
-------
service. Thus, 31 of the 41 valves per pump were
assumed to be in liquid service. The distribution
of these valves between light and heavy liquid service
was the same as that for the pumps.
The remaining valve distribution for hydrocarbon vapor
service, hydrogen service and light and heavy liquid
service was determined as follows:
1) Valves in liquid service = 31 x (number of pumps in
the process unit).
2) Valves in gas service = Total counted valves -valves
in liquid service (Step 1). It was assumed that
at least 10 percent of the valves in a process unit
are in gas stream service. The number of valves
calculated in this step can amount to less than 10
percent of the total valves. If so, the number of
gas valves is set at 10 percent of the total number
of valves.
3) Valves in hydrogen service = 30 percent of the total
number of valves in gas service for units which
utilize significant quantities of hydrogen (reforming,
HDS, hydrocracking).
4) Valve liquid distribution = Ratio of pump liquid
stream service as determined from this study.
Flanges - The number of flanges and other fittings was
determined to average 3.6 flanges per valve. This
factor was used to estimate the flange counts in
Table F4-3.
136
-------
— Compressors - The number of compressojrs was estimated
in a manner identical to that given for pumps. The
hydrocarbon-hydrogen categories are intended to dis-
tinguish between light hydrocarbon service compressors
and compressors in predominantly hydrogen service. The
number of seals (on which total emissions were based)
was determined to average two seals per compressor.
This value is given in parenthesis behind the actual
compressor count in Table F4-3.
Relief valves - From actual field counts, the number
of relief valves was determined to average six atmos-
pherically vented relief valves per process unit.
This factor was used to determine atmospheric emissions.
Drains - The number of drains per pump was determined
to average 2.6 in all counted process units. This factor
was used to estimate the drain counts in Table F4-3.
The source counts given in Table F4-3 refer to those
sources located within the battery limits of the process. That
is, equipment located in tankage or transfer lines is not
included in these source counts.
The number of sources per unit depends on a variety of
factors including processing complexity, process types, and
processing flexibility. Hence, the number of sources shown in
Table F4-3 is not necessarily representative of all or even the
majority of atmospheric distillation units.
In addition to the source counts and estimates devel-
oped by Radian, a second column of source counts is also given
in Table F4-3. These counts were based on information contained
137
-------
in Refejr-ence 8, The Development of Petroleum Refinery Plot Plans.
In this' study, process flow diagrams were used to determine the
number of pumps and compressors within the unit. The distribu-
tion of pumps between the light and heavy liquid service cate-
gories was based on pump service information obtained fron the
process flow sheets. The number of valves and fittings was based
on the number of pumps. For purposes of data comparison, the
valve counts listed in Reference 8 have been distributed between
the various service categories. This was accomplished using the
procedures described above for distributing the Radian valve
totals. In this case, however, the light and heavy liquid service
split was based on the pump distribution given in Reference 8.
The final column in Table F4-3 shows the estimated
fugitive emissions from each source category as well as an esti-
mate of emissions from the entire unit. Many of these estimates
are listed as a range, based on the two sets of source counts
described above.
It must be emphasized that all of the source counts
and stream service distributions given in Appendix B are, at
best, rough estimates. Even those values based on actual source
count data should be considered rough estimates since only a
small number of process units were counted. In addition, source
counts for similar types of process units showed large variations.
Therefore, reliable estimates for emissions source counts and
distributions should be obtained for the particular process unit
in question rather than using the estimates which are designed
to characterize typical refinery operation.
The estimated composition of fugitive non-methane
hydrocarbon emissions from an atmospheric distillation unit is
given in Table F4-4. These estimates indicate the types of
138
-------
TABLE F4-4.
U)
ESTIMATED COMPOSITION OF NON-METHANE HYDROCARBON FUGITIVE
EMISSIONS FROM A CRUDE DISTILLATION UNIT
Stream
Crude Straight run
Oil naphtha
Estimated percentage of emissions
attributed to each stream - %
Weighted contribution of each
component to unit emissions - pprnw
Benzene
Toluene
Ethylbenzene
Xylenes
Other Alkylbenzenea
Naphthalene
Anthracene
Blphenyl
Other Polynucleor aromatlca
n-Hexane
Other Alkanee
Oleflns
Cyclo Alkaneft
74
46
522
169
676
2871
660
108
246
6051
13820
673680
0
44770
24
59
617
208
382
3904
344
1
147
3528
9167
117660
0
99503
Middle
distillate
1
0
0
0
1
8
1
1
0
56
0
8627
0
1024
Atmospheric
gas oil
1
0
0
0
0
1
0
0
0
2
0
9724
0
512
Totals
100Z
105
1139
377
1059
6784
1005
110
393
9637
22987
809691
0
145809
999096
-------
hydrocarbons contributed by sources on a number-^of process streams
found within this unit. Additional information about these esti-
mates may be found in Appendix D.
4.1.2 Vacuum Distillation
Vacuum distillation is employed to fractionate topped
crude from the atmospheric distillation unit into a heavy resid-
ual oil and one or more heavy gas oil streams. The temperatures
necessary to distill the heavier crude fractions at atmospheric
pressure are so high that thermal cracking would occur. These
fractions are therefore distilled under vacuum to.prevent equip-
ment fouling and product loss resulting from cracking.
A vacuum distillation unit is an integral part of most
refineries. Charge capacity of U.S. vacuum distillation units
was more than 6.5 x 106 barrels per stream day in 1978.
4.1.2.1 Process Description--
A schematic diagram of a single-stage vacuum distilla-
tion unit is shown in Figure F4-2. Lubricant production some-
times necessitates adding a second stage to the primary vacuum
fractionator.
Distillation is carried out with absolute pressures in
the tower flash zone of 0.4-0.7 psi. In most units, the vacuum
inside the fractionator is maintained with steam ejectors,
although mechanical vacuum pumps may also be used. This system
is also designed to remove noncondensible hydrocarbon vapors
produced as the charge is heated.
140
-------
STEAM ?
HONCONDENSIBLE
GAS
TOPPED
CRUDE
SKAM
HEAVY VAC
GAS OIL
ASPHALT
70-1513-1
Figure F4-2. Vacuum distillation.
-------
7 Topped crude is heated in a direct-fired furnace and
charged to the vacuum fractionator. Vacuum distillation plus
steam stripping may be used to produce narrow boiling range lube
oil stocks for further processing. The injection of superheated
steam at the base of the column further reduces the partial
pressure of hydrocarbons in the fractionator to approximately 0.2
psi. Steam stripping is not required in the fractionation of
vacuum distillates for catalytic cracking or visbreaking feed-
stocks .
Distillation products are pumped from the fractionacion
tower at various points on the tower, with lighter components
distilling higher in the fractionator. Product specifications
depend on the crude feed and refinery design. Typical vacuum
distillation products include noncondensible refinery gas,
asphalt, and light and heavy vacuum gas oils. These gas oils may
be sent to a coker to be thermally cracked into gasoline feed-
stocks, or sent to a visbreaker to be cracked into distillate
fuel. Another alternative is hydrotreating, where sulfur can be
removed to upgrade the gas oils.
4.1.2.2 Process Technology
Process Conditions, Fuel, and Utility Requirements--
Typical operating parameters for a vacuum distillation unit are
listed below:
• Temperature-. 750-830°F
• Pressure: 0.4-0.7 psi
• Thermal energy: 74,900 Btu/bbl
142
-------
- -• • Electricity: 0.10-0.20 kW/bbl
• Steam: 8 Ib/bbl
Process Equipment--The major pieces of process equip-
ment in vacuum distillation units include heat exchangers, a
direct fired furnace, a vacuum fractionator, and a vacuum pro-
ducing system (usually steam ejectors). The vacuum fractionator
and steam ejectors are described briefly below.
Trays similar to th.ose used in atmospheric fractiona-
tors may also be used in vacuum fractionators. Tray efficiency
in vacuum columns tends to be much lower than that in atmospheric
columns, primarily because of the increased vapor loads and the
minimal liquid holding time on each tray. (Holding time must be
kept low to maintain low pressure drops across the trays.) Tray
efficiency may be increased by increasing the diameter of the
fractionation tower to a maximum of 40-45 feet.
Packed vacuum fractionation columns with diameters of
3-6 feet are replacing older, larger columns. These units con-
tain grid packing instead of trays, and operate with the liquid
phase trickling over the packing, countercurrent to the vapor
phase. A number of proprietary packing designs are in current
use.
Steam ejectors and barometric or surface condensers
are used in many vacuum distillation units to maintain operating
pressure. The size and number of ejectors and condensers used
is determined by the vacuum needed and the vapor load. To main-
tain a fractionator pressure of no more than 0.4 psia, three
ejector stages are usually required. The first stage condenses
the steam and compresses the fractionated noncondensible gases.
143
-------
The se"c<5nd and third stages remove the noncondensible gases from
the condensers.
The trend in vacuum distillation unit design is toward
the use of surface condensers rather than barometric condensers.
Surface condensers, although more costly than barometric conden-
sers, are much more effective in reducing hydrocarbon emissions
from the steam ejectors.
4.1.2.3 Atmospheric Emissions--
Emission sources from vacuum distillation units
include:
• Emissions from steam ejectors and barometric
condensers,
• Process heater flue gas emissions, and
• Fugitive emissions.
Vacuum System Emissions—Process hydrocarbon emissions
from steam vacuum ejectors have been estimated at 50 lb/103 bbl
charge. Using barometric condensers rather than surface conden-
sers may result in process emissions of as much as 1060 lb/103
bbl charge. Noncondensible hydrocarbon vapors removed by the
ejector system are released to the atmosphere unless combusted
in a furnace firebox or other type of combustion device.
Process Heater Emissions — Emission factors for various
pollutants from fired heaters are given in Table F4-5. A typical
vacuum distillation unit will require process heaters to raise
the temperature of the topped crude to between 750-830°F.
144
-------
TABLB-T4-5.
TYPICAL EMISSIONS FROM VACUUM DISTILLATION
UNIT PROCESS HEATERS
EPA Emission Factor
(lb/103 gal-oil fired)
(lb/106 scf-gas fired)
Total Emissions
(lb/103 bbl of
fresh feed)
Oil Fired Heaters
Participates
- Distillate oil
- Residual oil
1.1
Grade 4
Grade S
Grade 6
Sulfur Dioxide
- Distillate oil
- Residual oil
Sulfur Trioxidec
Carbon Monoxide
Hydrocarbons (as CH^)
Nitrogen Oxides
(as N02)
- Distillate oil
- Residual oilc
Gas Fired Heaters
Participates
Sulfur Oxides (as S02)
Carbon Monoxide
Hydrocarbons (as CHi,)
Nitrogen Oxides (as N02)
7
10
10(S)+3
142 (S)
157 (S)
2(S)
5
1
22
22+400 (N)
'
5-15
0.6
17
3
120-230
3.8
5.4
5.4(S)+1.6
76.1(S)
84.1(3)
l.KS)
2.7
0.54
12 ,
12+214 (N)
0.36-1.07
0.043
1.2
0.21
9.6-16.4
Source: Reference 9
Based on a heat input of 7.5X10*1 Btu/bbl of fresh feed with the following
fuel heating values: Oil - 140,000 Btu/gal; Gas - 1050 Btu/scf.
CS - Wt 2 sulfur in the fuel oil
Improper combustion may cause a significant increase in emissions
^se this emission factor for residual oils with less than 0.5% QJ<.5) nitro-
gen content. For oil with higher nitrogen content QJ>0,51, use emission
factor of 120 lb/103 gal
Based on sulfur content of 2000 gr/10G scf
145
-------
. ^.- Total emissions from vacuum unit heaters are also given
in Table F4-5. These figures are given as pounds of pollutant
per thousand barrels of feed using a typical heat input of 7.5 x
10 ** Btu per barrel of feed.
Fugitive Emissions — Fugitive emissions can occur from
a variety of sources including valves, pumps, compressors,
flanges and other fittings, relief valves, and drains. Hydro-
carbon emission rates for each individual source are quite low.
However, total fugitive emissions are significant because of the
large number of such sources within the refinery.
Total estimated fugitive non-methane hydrocarbon emis-
sions from a typical vacuum distillation unit are given in
Table F4-6. The listed emission factors were determined as a
part of this program. Additional information on these emission
factors, including a complete discussion on their derivation and
the confidence intervals for each source category, are contained
in Appendices B and C of this report.
Estimates for the number of sources within e&ch source
category were developed from:
• Actual field source counts conducted during the
course of this program, and
• Counts contained in The Development of Petroleum
Refinery Plot Plans, EPA publication number
450/3-78-025.8
The methods used to develop estimates for the number
of each source type are discussed in Section 4.1.1.3 of this
appendix.
146
-------
TABLE F4-6.
ESTIMATED FUGITIVE NON-METHANE HYDROCARBON EMISSIONS
FROM A TYPICAL VACUUM DISTILLATION UNIT
I 1
Emissions
Source Type
Valves
Open-End
(Sample)
Valves
i_i Pumps (Pump
-f> Seals)
-vl
Drains
Flanges &
Fittings
Relief Valves
Compressors
(Compressor
Seals)
aCounted From
Estimated
c Reference 8
rrocess
Stream Service
Clasfiif Ication
Gas /Vapor
Light Liquid
(V? > 0.\ pala 0 100'F)
Heavy Liquid
(VP < 0.1 pala 3 100'F)
Hydrogen Service
Total
All
Light Liquid
(VP > 0.1 pnla P 100'F)
Heavy Liquid
(VP < o.i pni« e IOO'P)
Total
All
All
All
Hydrocarbon
Hydrogen
Total
Flow Diagrams
Number of Sources in
Counts or Estimates
From Radian Study
50
45
405
0
~500b
-
2( 2)
14(20)
16(22)b
42b
U,
1785b
b
6
0
0
0
Process Unit
Councs or Estimates
From PES Studyc
71
142
497
0
710b
14»
2( 3)
7(10)
9(13)a
-
u
2350 b
-
0
0 .
0
, .
oource rtStlnmtec lot&l '(f'
Emission Emissions,
Factor, Ib/hr Ib/hr
0.059 2.95 - 4.19
0.024 1.08 - 3.41
0.0005 0.203- 0.249
0.018 0.0
4.23 - 7.85
0.005 0.070
0.25 0.50 - 0.75
0.046 0.46 - 0.92
0.96 - 1.67
0.070 2.94
0.00056 1.00 - 1.32
0.19 1.14
1.4 0.0
0.11 0.0
0.0 «
10.3 - 15.0
-------
i-~ The source counts given in Table F4-6 refer to those
sources located within the battery limits of the process. That
is, equipment located in tankage or transfer lines is not
included in these source counts.
The number of sources per unit depends on a variety of
factors including processing complexity, process type, and
processing flexibility. Hence, the number of sources shown in
Table F4-6 are not necessarily representative of all or even the
majority of vacuum distillation units.
4.1.3 Aromatics Extraction
The products of aromatic extraction—principally
benzene, toluene, and the xylenes (BTX)--are used in the petro-
chemical industry as intermediates in the production of plastics,
synthetic fibers, detergents, insecticides, synthetic rubber,
etc. Toluene is increasingly in demand as an ingredient in the
blending of gasoline to increase the octane rating of unleaded
gasoline. The extraction process is also used to produce
aromatic-free solvents for special applications.
Before the 1950s, the principal source of aromatic
hydrocarbons was coal tar. Now, virtually all aromatics are
extracted from two sources: the reformate produced by the
catalytic reforming of naphtha, and pyrolysis gasoline, produced
by the steam cracking of naphtha. Reformate has been for many
years the principal route to BTX in the U.S. Pyrolysis gasoline
is still important in Japan and, to a lesser extent, in western
Europe.
148
-------
4.1.3.1 i.- Process Description--
Aromatics extraction is a liquid/liquid solvent extrac-
tion process. The several proprietary processes are broadly
similar; the main difference between them is in the solvent that
is used.
Fresh feed, preheated to approximately 210°F, enters
the extraction column countercurrent to the lean solvent. Some
processes use rotating disks as contact stages; most use a series
of trays. Aromatics are selectively dissolved, and an almost
aromatic-free raffinate stream is withdrawn from the top of the
extractor. The raffinate is then cooled, and washed with water
to remove solvent.
The rich solvent leaves the bottom of the extractor and
is charged to a steam stripper. Some processes use a single
extractive stripping process to distill the extract into benzene,
toluene, and xylenes and to recover the solvent, but others
accomplish this in two stages. The stripper overhead is con-
densed into a stream containing some remaining aromatics and
saturated hydrocarbons. This stream is returned to the bottom
of the extractor column as reflux. The lean solvent is recycled.
Tetraethylene glycol, mixtures of several glycols,
dimethyl-sulfoxide, formal-morpholine, and tetrahydrothiphene-
dioxide are some of the solvents used.
There are a number of proprietary commercial extraction
processes. The Sulfolane and Udex processes account for the
majority of commercial installations for aromatics extraction:
each is in use in more than 50 refineries throughout the world.
The Tetra process is installed in more than 35 refineries.
149
-------
Sulfolan^-, originally developed by Royal Dutch/Shell, is licensed
by the UOP Process Division of UOP, Inc., as is Udex. The Tetra
licensor is Union Carbide Corporation. Most of the remaining
commercial installations are processes licensed by Howe-Baker
Engineers (Aromex), Snamprogetti S.p.A. (Formex), and the
Institut Frandais du Petrole (IFF).
4.1.3.2 Process Technology--
The aromatic content of a typical reformate is (vol. %)
benzene 5; toluene 24; ethylbenzene 4; para-, meta-, and ortho-
xylene 18; and C9 and Ci0 aromatics 4. The aromatic content of
typical pyrolysis gasoline is (weight %) benzene 32; toluene 14;
Ca aromatics 11; and €9+ aromatics 13. On the average, 100
percent of the benzene, 99 percent of the toluene, and 95-98
percent of the xylenes are recovered in the extraction process.
Fuel and utility requirements for the three most widely
used processes--Sulfolane, Udex, and Tetra--are given below.
Sulfolane Tetra
Utility per Bbl Feed Udex Process Process Process
Steam, Ib 400 2.5 125
Fuel, 103 Btu - 190
Cooling Water, gal . 1200 530 650
Electric Power, kWh 1.3 0.8 0.3
Operating parameters for the two most widely used
processes, Sulfolane and Udex, are given below.
150
-------
Sulfolane
Operating Condition Udex Process Process
Stripping steam ratio, wt/wt 0.6 0.13
Stripper bottom temperature, °F 290 375
Extractor top temperature, °F 290 212
Extractor pressure, psig 110 15
Feed temperature, °F 240 240
Process Equipment — The main equipment components of the
process are the extractor column, which contains trays or rotat-
ing disks to mix the feed and the solvent; one or more steam
stripping columns, which contain trays beneath which steam is
introduced to contact steam and rich solvent; and a conventional
fractional distillation unit where the extract is separated into
benzene, toluene, and xylenes.
Depending on the particular proprietary process, a
water still and a vacuum solvent regenerator may also be required,
4.1.3.3 Atmospheric Emissions--
Since aromatics extraction is a closed process, the
only significant emissions are fugitive hydrocarbon emissions.
Fugitive Emissions — Fugitive emissions can occur from
a variety of sources including valves, pumps, compressors,
flanges and other fittings, relief valves, and drains. Hydro-
carbon emission rates for each individual source are quite low.
151
-------
However^; total fugitive emissions are significant because of the
large number of such sources within the refinery.
Total estimated fugitive non-methane hydrocarbon emis-
sions from a typical aromatics extraction unit are given in
Table F4-7. The listed emission factors were determined as a
part of this program. Additional information on these emission
factors, including a complete discussion on their derivation and
the confidence intervals for each source category, are contained
in Appendices B and C of this report.
Estimates for the number of sources within each source
category were developed from:
• Actual field source counts conducted during the
course of this program, and
• Counts contained in The Development of Petroleum
Refinery Plot Plans, EPA publication number
450/3-78-025. 8
The methods used to develop estimates for the number
of each source type are discussed in Section 4.1.1.3 of this
appendix.
The source counts given in Table F4-7 refer to those
sources located within the battery limits of the process. That
is, equipment located in tankage or transfer lines is not
included in these source counts.
The number of sources per unit depends on a variety of
factors including processing complexity, process type, and
processing flexibility. Hence, the number of sources shown in
152
-------
TABLE F4-7.
LO
ESTIMATED FUGITIVE NON-METHANE HYDROCARBON EMISSIONS
FROM A TYPICAL AROMATIC EXTRACTION UNIT
Emissions
Source Type
Valves
Open-End
(Sample)
Valves
Pumps (Pump
Seals)
Drains
Flanges &
Fittings
Relief Valves
Compressors
(Compressor
Seals)
Process
Stream Service
Classification
Gas/Vapor
Light Liquid
(VP > 0.1 pain 9 100'P)
Heavy Liquid
(VP < 0.1 psla « 100'P)
Hydrogen Service
Total
All
Llglit Liquid
(VP > 0.1 pala S 100'P)
Heavy Liquid
(VP < 0.1 pala S 100° f)
Total
All
All
All
Hydrocarbon
Hydrogen
Total
Number of Sources In
Counts or Estimates
From Radian Study
60
486
54
0
600b
-
16(23)
2( 3)
18(25)b
2142b
6h
0
0
0
Process Unit
Counts or Estimates
From PES Study
206
1370
483
0
2059b
29a
17(24)
6( 8)
23(32)
-
6815b
-
0
0
0
Source
Emission
Factor, Ib/hr
0.059
0 . 024
0.0005
0.018
0.005
0.25
0.046
0.070
0.00056
0.19
1.4
0.11
Estimated Total
Emissions, , , ,.
Ib/hr
3.54 - 12.2
11.7 - 32.9
0.027 - 0.242
0.0
15.3 - 45.3
0.145
• 5.75 - 6.00
0.138 - 0.368
5.89 - 6.37
3.29
1.20 - 3.82
1.14
0.0
0.0
0.0
27.0 - 59.6
Counted From Flow Diagrams
Estimated
p
Reference 8
-------
Table, f^-7 are not necessarily representative of all or even the
majority of aromatics extraction units.
The estimated composition of fugitive non-methane
hydrocarbon emissions from an aromatics extraction unit is given
in Table F4-8. These estimates indicate the types of hydrocar-
bons contributed by sources on a number of process streams found
within this unit. Additional information about these estimates
may be found in Section 2.4 of Appendix D.
4.1.4' References for Section 4.1
I. Refining Process Handbook. Hydrocarbon Processing, 57(9),
1978.
2. Hobson, G. D., and W. Pohl. Modern Petroleum Technology.
Applied Science Publishers Ltd., Essex, Great Britain, 1975.
3. Good Vaccum Unit Design Pays Off. Oil and Gas Journal, 76
(12): 72-79, March 13, 1978.
4. Annual Refining Survey. Oil and Gas Journal, 77(13): 122-
127, March 26, 1979.
5. Gary, J. H., and G. E. Handwerk. Petroleum Refining: Tech-
nology and Economics. Manual Dekker, Inc., New York, 1975.
pp. 31-46.
6. Dickerman, J. C., et al. Industrial Process Profiles for
Environmental Use: Chapter 3. Petroleum Refining Industry.
EPA-600/2-77-023c. U.S. Environmental Protection Agency,
Washington, D.C., January 1977.
7. Wetherold, R. G. The Distribution of Selected Fugitive
Hydrocarbon Emission Sources Among Petroleum Refinery Pro-
cess Streams. EPA Contract No. 68-0202665. U.S. Environ-
mental Protection Agency, Research Triangle Park, N.C, 1979,
unpublished.
8. U.S. Environmental Protection Agency, Office of Air Quality
Planning and Standards. Development of Petroleum Refinery
Plot Plans. EPA-450/3-78-025. Research Triangle Park,
North Carolina, 1978.
154
-------
TABLE F4-8.
Ln
ESTIMATED COMPOSITION OF FUGITIVE NON-METHANE HYDROCARBON
EMISSIONS FROM AN AROMATICS EXTRACTION UNIT
Stream
Estimated percentage of emissions
attributed to each stream - %
Weighted contribution of each
component to unit emissions - ppmw
Benzene
Toluene
Ethylbenzene
Xylcncs
Other Alkylbenzenes
Naphthalene
Anthracene
Blphenyl
Other Polynuclear aromatlcs
n-Hexane
Other Alkanea
Oleflns
Cyclo Alkanes
Re formate
12
648
9324
4020
20508
38928
858
0
0
84
2880
42720
0
0
Solvent
0
0
0
0
0
o •
0
0
0
0
0
0
0
0
Aromatic
Extract
44
7850
112948
48695
248420
21120
44
0
0
44
44
836
0
0
Rafflnate
44
22
330
132
660
1012
22
0
0
22
27720
410080
0
0
Totals
100%
8520
122602
,52847
269588
61060
954
0
0
150
30644
453637
0
0 »
1000000
-------
9. Environmental Protection Agency. Compilation of Air Pollu-
tant Emission Factors. Third Edition, Supplement No. 8.
AP-42. Research Triangle Park, NC, May 1978.
10. Air Pollution Control District, County of Los Angeles; State
of California, Department of Public Health; U.S. Department
of Health, Education and Welfare; Western Oil and Gas
Association. Emissions to the Atmosphere from Eight Miscel-
laneous Sources in Oil Refineries. June 1958.
11. Deal, G. H., Jr., et al., A Better Way to Extract Aromatics.
Petroleum Refiner, September 1959, p. 195.
156
-------
This coke is then recovered and sold as a by-product; its value
is dependent on the quality of the feedstock.1
Delayed coking is likely to remain an important
refining process since it simultaneously converts low value
materials to lighter, more valuable materials while producing
coke as a valuable by-product. More detailed information on
this process is contained in Section 4.2.1.
Fluid coking is a continuous process which converts
heavy feedstocks such as residuum and pitch into lighter, more
valuable products. The heated feed is sprayed into a fluidized
bed of hot coke particles causing additional coke to form. The
cracked oil vapors leave the reactor and are fractionated into
a heavy recycle stream, various gas oil streams, and gasoline.
The coke particles are sent to a burner where they are heated by
burning a portion of the coke with a limited amount of air. The
hot coke particles are then recirculated back to the reactor
where they furnish much of the heat required for the cracking
reactions. Since there is a net production of coke in the unit,
coke is removed from the burner as a by-product.1
Unlike delayed coke, fluid coke is often unsuitable
for many end uses, and is consequently less valuable than
delayed coke. Additional information on fluid coking is pre-
sented in Section 4.2.2.
Visbreaking is a relatively mild cracking operation
used to reduce the viscosity of straight-run residues such as
pitch. These heavy residues are often blended with lighter
heating oil's to produce fuel oils of acceptable viscosity. This
procedure, however, uses a high value product, light heating
oil, to produce a lower value product, fuel oil. By reducing
158
-------
the viscosity of the pitch fraction, visbreaking reduces the
amount of required light heating oil.1
The visbreaking reactions occur within specially
designed tubes in a fired heater. The heated product is then
fractionated into low quality gasoline, light and heavy gas oil,
and a residual tar of reduced viscosity. Visbreaking is dis-
cussed in further detail in Section 4.2.3.
Various thermal cracking operations have been used in
petroleum refining since 1910. However, thermal cracking of
high boiling materials is no longer a major source of motor
gasoline since its octane rating is low.
A few thermal cracking units are still in operation.
These processes utilize temperatures of 850-1000°F and pressures
ranging from 100-1000 psig. Products include light hydrocarbon
gases, cracked gasoline, light heating oil, and thermal tar.
Thermal tar can be used as coker feedstock or it can be blended
into fuel oil.
The current (1979) utilization by U.S. refiners of the
processes described above is supplied below.3
Process
Delayed coking
Fluid coking
Visbreaking
Thermal cracking (gas oil
or reduced crude feed)
Number of Units
48
4
13
15
Total Capacity,
bbl/sd
997,900
77,000
160,500
150,600
159
-------
4.2.1- r' Delayed Coking
The delayed coking process was developed as a method
to reduce the yield of residual fuel oil by severe thermal
cracking of feeds such as atmospheric or vacuum residuals, and
thermal tars. In addition, delayed cokers are capable of pro-
cessing other materials including full range crude oils, shale
oil, bitumin, gilsonite, and coal tar pitch.
Products from delayed coking consist of light gases,
gasoline, diesel, light and heavy gas oil, and coke. Gas oils,
which constitute the largest portion of the liquid products, are
typically utilized as catalytic cracking feedstock for the
production of high octane gasoline. Technological improvements
in downstream processing now permit the economical upgrading of
other coker products. For example, coker gasoline can be
catalytically reformed to high octane gasoline after suitable
hydrotreating for the removal of metals, nitrogen, or sulfur
contaminants.
Solid petroleum coke is produced as a product in
delayed coking operations. The main uses of petroleum coke
include:2
• Manufacture of electrodes for the electric
furnace production of elemental phosphorus,
titanium dioxide, steel, calcium carbide,
and silicon carbide.
• Manufacture of anodes for electrolytic cell
reduction of alumina.
160
-------
- i.- • Direct use as carbon source for production
of phosphorus, calcium carbide, and silicon
carbide.
• Manufacture of graphite.
• Direct use as fuel.
The quality of the coke produced in delayed coking
depends in part on the operating conditions and the quality of
the feed. Desirable qualities in petroleum coke include low
porosity, a low coefficient of thermal expansion, good electri-
cal conductivity, and low ash, metals, and sulfur content.
The most common form of coke, is produced as hard,
irregular-shaped lumps called sponge coke. Sponge coke is
generally produced when using low quality feeds such as vacuum
residuals with high asphaltene content. During the coking pro-
cess, the asphaltenes, which exist as colloidal suspensions,
form a highly cross-linked amorphous coke with a high concentra-
tion of impurities. These impurities, which include sulfur,
nickel, and vanadium, make sponge coke unsuitable for high
grade anodes.6
Some highly aromatic materials, such as catalytic
cracker cycle oil and thermal cracking tars, are used to produce
high quality "needle" coke. Needle coke derives its name from
its microscopic elongated crystalline structure. This structure
is a result of condensation and polymerization of large aromatic
molecules. Since needle coke is produced from previously
cracked oils which contain less sulfur than residues, the sulfur
content of needle coke is generally lower than that of sponge
coke. An additional reason for lower needle coke sulfur content
161
-------
could-h£- the differences in sulfur distribution, between the
aromatic and asphaltene molecules. That is, sulfur is not
readily released from the interior of asphaltene molecules which
form sponge coke. Needle coke is preferred over sponge coke for
electrode manufacture because of its lower electrical resis-
tivity and its lower coefficient of thermal expansion.5'6
4.2.1.1 Process Description
A simplified flow diagram for the delayed coking
process is given in Figure F4-3. The process flow is quite
simple and has remained relatively unchanged since the process
was developed. Fresh feed is fed directly to the bottom of the
main fractionator where materials lighter than the desired end
point of the heavy gas oil product are flashed off. The
remaining material combines with the recycle and is pumped to
the coking heater where the temperature of the oil is quickly
raised to around 950°F. The liquid-vapor mixture leaving the
coking heater passes to a coke drum where the coking reactions
occur. The coke drum is actually just a wide spot in the line
where sufficient residence time is provided for the reactions
to go to completion.
Most coking units use two drums (or two sets of drums).
After the first drum becomes filled with coke, the flow of hot oil
is diverted to the second drum. Thus, a semicontinuous process-
ing scheme is established by alternating the coking and decoking
operations for each drum. The coke drums are often sized so
that each one operates on a 48-hour cycle, permitting the de-
coking operation to be scheduled at the same time each day on a
24-hour cycle.
162
-------
JT'I
GrtS
COKE
DRUMS
HEATER
1 930'F
, i CONDENSAIC
J f DRUM
Q D
COKE
FEtO
Figure F4-3. Process flow diagram - delayed coking.
Source: Reference 5
-------
i.~ Hot vapor from the coke drum returns -to the
fractionator for product separation. The heaviest portion
combines with fresh feed in the fractionator and is recycled
back to the heater.
The design of the fractionator above the gas oil tray
is similar to that employed for crude towers. Below that point,
however, special designs are required. As mentioned previously,
fresh feed is introduced near the bottom of the tower, perhaps
2 to 4 trays above the bottom vapor space. This accomplishes
the following:2
• Hot vapors from the coke drum are quenched
preventing significant coke formation
within the fractionator.
• A portion of the hot coke drum vapors are
condensed for recycle.
• Any material lighter than the desired coke
drum feed is stripped from the fresh feed.
• The fresh feed liquid is further preheated.
The design of the coker heater is also important to
the economical operation of the unit. Since the mixture is
heated to temperatures above the coking temperature, steps must
be taken to suppress coke buildup in the heater tubes. Gener-
ally, coke formation is suppressed by passing the oil through
the heater at high velocity. As the oil nears the coking
temperature, an increasing quantity of the feed is vaporized,
serving to increase the velocity and turbulence even further.
164
-------
If insufficient velocity is obtained by vaporisation of the oil,
steam or water can be injected. With sufficient precautions,
coker heaters can operate for periods of well over a year before
decoking of the heater tubes is required.0'2
Decoking the coke drums consists of numerous individ-
ual operations. After switching drums, steam is introduced to
the bed for removal of volatile combustion matter (VCM). Then,
water is injected at a controlled rate to cool the coke. After
the excess water is drained, the coke drum is opened and the
coke removed.
The first step in the decoking operation is to drill
a pilot hole through the bed. This pilot hole.is then enlarged
to permit lumps of coke to fall through the bed without the
danger of coke buildup around the drill stem. The coke is then
cut away from the drum using either a mechanical drill or, more
commonly, a hydraulic system. The hydraulic system consists of
a number of high pressure (2000 to 3000 psig) water jets which
are lowered into the bed on a rotating stem. The coke is then
cut from the drum, a layer at a time, starting from either the
top or the bottom.5
Several methods of collecting the coke are available.
In many cases, the coker is mounted over a railroad track so
that the coke can be discharged directly into railroad cars.
The coke is retained in the cars while water and coke fines
drain off and flow to a sump. Alternatively, the coke can be
directed to a concrete apron or pit. The water and coke fines
drain off, and the coke is moved using a scoop or bridge crane.
In either system, water collected in the sump is usually sent
to a clarifier for fines removal. Clarified water is then sent
to a storage tank for reuse.
165
-------
7 After all coke has been removed, the "drum is closed,
pressure tested, and reheated at a controlled rate. The entire
decoking operation is completed in 24 hours using a schedule
similar to that given below. 2 ' 5
Operation Time, hours
Fill drum with coke 24.0
Switch drums - 0.5
Steam Out 3 .0
Cooling 3.0
Drain '. 2.0
Unhead and Decoke - 5.0-
Head up test 2.0
Warm-up 7.0
Spare time 1.5
Total 48.0
4.2.1.2 Process Technology--
Processing Variables — Important process variables for
delayed coking include the heater outlet temperature, the
fractionator pressure, the vapor temperature at the heavy gas
oil draw-off tray, and the Conradson "free" carbon content of
the feed. At higher heater outlet temperatures, the rate of the
cracking and coking reactions increases. This increases the
yield of gas, naphtha, and coke at the expense of a lower gas
oil yield. Typical heater outlet temperatures range from 900-
975°F.3
166
-------
~ Increasing the fractionator pressure "has much the same
effect as an increase in temperature. Higher pressure in the
fractionator serves to condense more recycle within the frac-
tionator for return to the heater and coke drums. Thus, less
gas oil is produced relative to the amount of gas, naphtha, and
coke. Typical pressures used in delayed coking range from 20-60
psig. However, when high quality aromatic feedstocks are used,
the pressure can be increased to around 100 psig in the coke
drum to promote the formation and yield of valuable needle
coke.
The temperature of the vapor at the gas oil draw-off
tray determines the end point of the gas oil product. If this
temperature is raised, less gas oil is available for recycle.
Thus, the yield of gas oil is increased at the expense of gas,
naphtha, and coke. Often, the gas oil end point temperature
is dictated by the allowable metals content.
The Conradson carbon residue (CCR) is a laboratory
measurement of the coke forming potential of a particular stock.
And, a number of correlations are available which relate CCR
to the expected commercial coke yield.5
The recycle ratio is often given as an operating
variable in coking operations. Actually, the recycle ratio is
a function of the fractionator pressure and the gas oil draw-
off tray vapor temperature. High recycle ratios are used when
processing for maximum gas oil. Typical recycle ratios range
from 0.1-1.0.2
167
-------
The following is a summary of operating conditions
for delayed coking:l'2'"'5
Heater Outlet Temperature, °F 900-975
Coke Drum Pressure, psig 20-100
Recycle Ratio 0.1-1.0
4.2.1.3 Process Emissions--
Emission sources from delayed coking operations
include:
• Emissions during decoking operations,
• Process heater flue gas, and
• Fugitive emissions.
Emissions During Decoking Operations—At the end of
each coking cycle, coke is removed from the coke drum. This
procedure has been described in a previous section (4.2.1.1).
Part of the cooling operation involves injecting steam into the
coke drum. The procedure serves to remove the majority of the
hydrocarbons remaining in the drum. The steam is then condensed
and the remaining vapors are usually flared.
Following the cool-down procedure, the drum is opened
and the coke is removed. This operation is a source of particu-
late emissions to the atmosphere. In addition, hydrocarbon
vapors entrained in the coke may be released during the cutting
operation. A water quench is often used to minimize particulate
emissions. Since this water will contain some sulfur compounds,
163
-------
it may ire a source of objectionable odors. Emission factors for
the decoking operation are unavailable.
Process Heater Emissions — Emission factors for various
pollutants from fired heaters are given in Table F4-9. A
typical delayed coking unit will require process heaters to
raise the temperature of the coke drum feed to the desired
coking temperature.
Total emissions from the delayed coking heaters are
also given in Table F4-9. These figures are given as pounds
of pollutant per thousand barrels of feed using a typical heat
input of 2.4 x 105 Btu per barrel of feed.11*
Fugitive Emissions — Fugitive emissions can occur from
a variety of sources including valves, pumps, compressors,
flanges and other fittings, relief valves, and drains. Hydro-
carbon emission rates for each individual source are quite low.
However, total fugitive emissions are significant because of the
large number of such sources within the refinery.
Total estimated fugitive non-methane hydrocarbon
emissions from a typical delayed cooking unit are given in Table
F4-10. The listed emission factors were determined as a part of
this program. Additional information on these emission factors,
including a complete discussion on their derivation and the
confidence intervals for each source category, are contained
in Appendices B and C of this report.
Estimates for the number of sources within each source
category were developed from:
169
-------
?ABLE F4-9.
TYPICAL EMISSIONS FROM DELAYED COKING
UNIT PROCESS HEATERS
EPA Emission Factor
(lb/103 gal-oil fired)
(lb/106 scf-gas fired)
Total Emissions
(lb/103 bbl of
coker feed)
Oil Fired Heaters
Particulates
- Distillate oil
- Residual oil
Grade 4
Grade 5
Grade 6
Sulfur Dioxide0
- Distillate oil
- Residual oil
Sulfur Trioxidec
Carbon Monoxide
Hydrocarbons (as
Nitrogen Oxides
(as N02)
- Distillate oil
- Residual oil6
7
10
10(S) + 3
142(S)
157(S)
2(S)
5
1
22
22+400(N)
3.4
12
17
17(S) + 5.1
243 (S)
269(S)
3.4(S)
8.6
1.7
38
38+686(N)"
Gas
Fired Heaters
Particulates
Sulfur Oxides (as S02)f
Carbon Monoxide
Hydrocarbons (as CEj,)
Nitrogen Oxides (as N02)
5-15
0.6
17
3
120-230
1.1-3.4
0.14
3.9
0.69
27-53
*Source: Reference 15
Based on a heat input of 2.4xlOs Btu/bbl of fresh feed with the following
fuel heating values: Oil - 140,000 3tu/gal; Gas - 1050 Btu/scf.
CS - Wt % sulfur in the fuel oil
Improper combustion nay cause a significant increase in emissions
^se this emission factor for residual oils with less than 0.52 CN-s.5) nitro-
gen content. For oil with higher nitrogen content (N>0,5), use emission
factor of 120 lb/103 gal
Based on sulfur content of 2000 gr/106 scf
170
-------
TABLE F4-10.
ESTIMATED FUGITIVE NON-METHANE HYDROCARBON EMISSIONS
FROM A TYPICAL DELAYED COKING UNIT
Emissions
Source Type
Valves
Open-End
(Sample)
Valves
Process
Stream Service
Classification
Gas/Vapor
Light Liquid
(VP > 0.1 pala 0.1 pala 9 100'F)
(VP
Drains
Flanges &
Fittings
Relief Valves
Compressors
(Compressor
Seals)
Heavy Liquid
< 0.1 pala 9 100'P)
Total
All
All
All
Hydrocarbon
Hydrogen
Total
2( 3)
7(10).
9(13)b
23b
K
1071°
6b
0
0
0
18(25)
4( 6)
22(31)a
-
K
5875°
-
3( 6)
0( 0)
3( 6)
0.25
0.046
0.070
0.00056
0.19
1.4
0.11
0.750
0.460
1.21
0.60
0.0
0.0
8.03
- 6.25
- 0.276
- 6.53
1.61
- 3.29
1.14
- 8.4
0.0
- 8.4
- 63.5
Counted From Flow Dlagrama
Estimated
Reference 14
-------
7 • Actual field source counts conduc-ted during
the course of this program, and
• Counts contained in The Development of Petroleum
Refinery Plot Plans, EPA publication number
450/3-78-025.llf
The methods used to develop estimates for the number of
each source type are discussed in Section A. 1.1.3 of this
appendix.
The source counts given in Table F4-10 refer to those
sources located within the battery limits of the process. That
is, equipment located in tankage or transfer lines is not included
in these source counts.
The number of sources per unit depends on a variety of
factors including processing complexity, process type, and pro-
cessing flexibility. Hence, the number of sources shown in
Table F4-10 are not necessarily representative of all or even
the majority of fluid coking units.
The estimated composition of fugitive non-methane
hydrocarbon emissions from a delayed coking unit is given in
Table F4-11. These estimates indicate the types of hydrocarbons
contributed by sources on a number of streams within the process
unit. Additional information about these estimates may be found
in Section 2.4 of Appendix D.
172
-------
TABLE FA-11.
ESTIMATED COMPOSITION OF FUGITIVE NON-METHANE HYDROCARBON
EMISSIONS FROM A DELAYED COKING UNIT
-vj
U)
Scream
Estimated percentage of emissions
attributed Co each stream - 1
Weighted contribution of each
component to unit emissions — ppraw
Benzene
Toluene
Ethylbenzene
Xylenes
Other Alkylbenzenes
Naphthalene
Anthracene
Blphenyl
Other Polynuclear aromatico
n-llexane
Other Alkanen
Oleflns
Cyclo Alkanes
Hydrogen
Vacuum
Resld
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
Coke
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
Cracked
Naphtha
57
1642
51175
12215
97727
138778
6242
0
0
3694
6743
116340
97322
38122
0
LPG
Oleflns
14
0
0
0
0
0
0
0
0
0
0
56000
84000
0
0
Fuel
Gas
29
0
0
0
0
0
0
0
0
0
0
266800
17400
0
5800
Totals
100X
1642
51175
12215
97727
138778
6242
0
0
3694
6743
439140
198722
38122
5800
1000000
-------
4.2.2- i.- Fluid Coking
Fluid coking, like delayed coking, was developed as a
method of reducing the overall yield of residuals, tars, and
resins produced during certain refining operations. These
materials are converted to lighter, more valuable liquid products
and coke.
4.2.2.1 Process Description--
The yield of liquid products from fluid coking is
similar to that of delayed coking. Coke production, however,
is significantly lower in fluid coking. Factors leading to
decreased coke production include:
• Higher yields of heavy gas oil, and
• A substantial portion of the coke is
consumed within the unit to provide process
heat.
The coke produced by the fluid coking process is
usually of insufficient quality for most industrial uses. Con-
sequently, most of this coke is sold as fuel or used within the
refinery to produce steam.1'10 The fact that many delayed
coking units are now being operated to produce valuable premium
cokes may account for the limited use of fluid coking process.
In fact, the number of fluid coking units in operation has
actually decreased from 10 units in 1970 to just 4 units in
1979.3
A relatively recent advancement in coking processes
is Flex_icoking, developed by the Exxon Research and Engineering
174
-------
Company.-. Flexicoking integrates conventional .fluid coking with
coke gasification to produce a coke gas containing significant
quantities of carbon monoxide, hydrogen, carbon dioxide, water
vapor, and nitrogen. The coke gas may be substituted for
refinery fuel gas or natural gas to fire process heaters or
boilers.9 The net coke production from Flexicoking is reported
to be in the range of 1-2 wt. percent of the feed as compared to
10-20 percent for conventional fluid coking. Up to 99 percent
of the metals are concentrated in the coke by-product. At this
time no commercial Flexicokers have been installed in the United
States.9 Extensive commercial experience has been accrued
abroad, however, mainly in Japan.
4.2.2.2 Process Technology--
As indicated in the simplified process flow diagram
given in Figure F4-4, fluid coker units contain two major
process vessels: a fluidized bed reactor and a coke burner.
Coke particles are continuously circulated between these two
vessels to provide process heat. The oil or pitch feed is
injected directly into the bed of hot coke particles. Part of
the feed is vaporized while the remainder is distributed on the
surface of the coke where the cracking reactions occur. The
cracked products vaporize leaving a thin layer of new coke on
the particles. Steam, injected into the bottom of the bed,
keeps the bed fluidized and promotes mixing. This steam also
serves to strip any remaining hydrocarbons from the coke par-
ticles as they leave the reactor bed.
The vapor products are passed through cyclones and
charged to the bottom of a scrubber. The scrubber serves to
condense the heaviest components which are recycled to the
reactor. The vapors leaving che scrubber are sent to a
fractionator for product separation.
175
-------
To
Condenser •*
Gas Oil
Recycle •*
Flue Gas
Steam
Water
Product Coke
to Storage
Figure F4-4. Process flow diagram - fluid coking.
Source: Reference 14
-------
- i.- The steam-stripped coke particles leaving the reactor
proceed to the burner vessel. Within the burner, approximately
25 percent of the coke formed in the reactor is consumed.8 The
heat produced by combustion heats the remaining coke particles
which are then returned to the reactor. No preheat furnace for
the feed is required since the circulating hot coke particles
provide the heat necessary for the coking reactions.
Since more coke is produced in the reactor than is
consumed in the burner, a stream of coke is withdrawn from the
burner. The coke is water-cooled in a quench elutriator and
sent to storage.
The flue gas from the burner bed normally contains
significant quantities of carbon monoxide and has a heating value
of around 43 Btu/scf.n Hence, this stream is usually discharged
through cyclones to a waste heat recovery device such as a CO
boiler or process heater.
A simplified flow diagram for the Flexicoking process
is given in Figure F4-5. This process is a combination of con-
ventional fluid coking and coke gasification. The Flexicoking
process utilizes three process vessels; a reactor, a coke
burner, and a coke gasifier. The operation of the reactor and
burner vessels is nearly identical to that of conventional
fluid coking. However, most of the coke produced within the
unit is sent from the burner to the gasifier. The gasification
is similar to fluid bed coal gasification practiced commercially
since the 1920's in the Winkler process. Basically, steam and
air are reacted with the coke at elevated temperatures to
produce a coke gas with a heating value of 100-130 Btu/scf.9
177
-------
COKE GAS
—i
00
REACTOR PRODUCTS
TO" FRACTIONATOR
SCRUBBER
I T
RECYCLE FEED
r
RESID
FEED
r
REACTOR
STEAM
STEAM
GENERATION
HEATER
COOLING
T
FINES
REMOVAL
SULFUR
REMOVAL
SULFUR
COKE FINES
GASIFIER
BLOWER
.AIR
STEAM
Figure F4-5. Process flow diagram - Flexicoking.
Source: Reference L2
-------
- i-' The coke gas is returned to the burner vessel where
it is cooled to provide a portion of the heat carried back to
the reactor. The combined coke gas and burner flue gas stream
leaves the burner vessel where it is cooled and desulfurized.9
Approximately 95 percent of the total coke produced is
either consumed in the burner or converted to coke gas in the
gasifier. The remaining 5 percent (%l-2 wt. percent of the
fresh feed) is removed from the burner. This coke has been
substantially desulfurized but contains up to 99 percent of the
metals present in the feed.9.
Fluid coking and Flexicoking operating conditions are
summarized in Table F4-12.
TABLE F4-12.
PROCESS CONDITIONS FOR FLUID
COKING AND FLEXICOKING
Fluid Coking
Flexicoking
Temperature, °F
Reactor
Burner
Gasifier
950
1150
950
1150
1300-1800
Pressure, psig
Reactor
Burner
Gasifier
10
11
10
11
25-45
Utilities - per barrel feed
Electric Power, kWh 5.5
Fuel, MM Btu 0
13
0
Source^ References 1, 4, 8 and 9
179
-------
4.2.2-. 2-' Process Emissions-- -
Emission sources for the fluid coking process include:
• The burner vessel flue gas
• Fugitive emissions.
Operating fluid cokers do not require the use of a
process heater since heat is obtained from burning coke produced
by the unit.
Burner Vessel Emissions--
Approximately 25 percent of the coke formed within the
fluid coker reactor is consumed within the burner vessel. And,
the heat generated by combustion heats the remaining coke
particles which are then returned to the reactor. The combus-
tion process produces a flue gas containing substantial quanti-
ties of carbon monoxide with lesser amounts of SOX, NOX,
organics, and particulates.
CO emission rates have been estimated at 30 pounds per
barrel of feed.13 And, venting of this gas would represent a
substantial loss of energy. At present, all four domestic fluid
coking units utilize either a CO boiler or a fired heater to
recover this energy.
Use of a CO boiler will reduce the concentration of CO
and other combustible materials to low levels. There may also
be a reduction in particulate emissions. However, high combus-
tion temperatures may cause an increase in NOX emissions and
both SO^, and NOX will increase if auxiliary fuel is required.
180
-------
~ Fugitive Emissions--Fugitive emissions can occur from
a variety of sources including valves, pumps, compressors,
flanges and other fittings, relief valves, and drains. Hydro-
carbon emission rates for each individual source are quite low.
However, total fugitive emissions are significant because of the
large number of such sources within the refinery.
Total estimated fugitive non-methane hydrocarbon emis-
sions from a typical fluid coking unit are given in Table F4-13.
The listed emission factors were determined as a part of this
program. Additional information on these emission factors, in-
cluding a complete discussion on their derivation and the con-
fidence intervals for each source category, are contained in
Appendices B and C of this report.
Estimates for the number of sources within each source
category were developed from:
• Actual field source counts conducted during the
course of this program, and
• Counts contained in The Development of Petroleum
Refinery Plot Plans, EPA publication number
450/3-78-025.1*
The methods used to develop estimates for the number
of each source type are discussed in Section 4.1.1.3 of this
appendix.
The source counts given in Table F4-13 refer to those
sources located within the battery limits of the process. That
is, equipment located in tankage or transfer lines is not
included in these source counts.
181
-------
TABLE F4-13.
oo
N)
ESTIMATED FUGITIVE NON-METIIANE HYDROCARBON EMISSIONS
FROM A TYPICAL FLUID COKING UNIT
Emissions
Source Type
Valves
Open-End
(Sample)
Valves
Pumps (Pump
Seals)
Drains
Flanges &
Fittings
Relief Valves
Compressors
(Compressor
Seals)
Process
Stream Service
Classification
Gas/Vapor
Light Liquid
(VP > 0.1 pala 8 100'F)
Heavy Liquid
(VP < o.i p«i« « loo'p)
Hydrogen Service
Total
All
Light Liquid
(VP > 0.1 puts 9 100'P)
Heavy Liquid
(VP < 0.1 pala 9 100'F)
Total
All
All
All
Hydrocarbon
Hydrogen
Total
Number of Sources In Process Unit
30
58
216
0
304a
—
2( 3)
7(10)
9(13)a
28a
1047a
6b
4(8)
0(0)
4(8)*
Source
Emission
Factor, Ib/hr
0.059
0.024
0.0005
0.018
0.005
0.25
0.046
0.070
0.00056
0.19
1.4
0.11
Estimated Total
Emissions.
Ib/hr ^
1.77
1.39
0.108
0.0
3.27
~
0.75
0.46
1.21
1.96
0.586
l.U
11.2
0.0
11.2
19.4 ,|
Physically Counted
^Estimated
-------
^.- The number of sources per unit depends- on a variety of
factors including processing complexity, process type, and pro-
cessing flexibility. Hence, the number of sources shown in
Table F4-13 are not necessarily representative of all or even the
majority of fluid coking units.
4.2.3 Visbreaking
Viscosity breaking or visbreaking is a mild thermal
cracking operation used to reduce the viscosity of materials
such as atmospheric or vacuum residuals and pitch. This pro-
cedure reduces the amount of valuable light heating oil which
must be blended with the residuum to produce a fuel oil of
acceptable viscosity. Currently, 13 visbreaking units are
in operation with a total capacity of 160,500 bbl/day.3
4.2.3.1 Process Description and Technology--
A flow diagram for the visbreaking process is given
in Figure F4-6. The reduced crude charge is heated in a fired
heater where mild liquid phase cracking occurs. The heater
effluent is quenched and then flashed in a combination tower.
The flashed vapors can be fractionated into gas, low quality
gasoline, light and heavy gas oil, and a visbreaker tar of
reduced viscosity. Alternatively, the gas oil fraction can be
included with the bottoms product which can further reduce the
amount of lighter cutter stock required to meet viscosity
specifications.*
Visbreaking processes utilize temperatures at the
heater outlet of 850-950°F. And, the pressure at this point is
in the range of 50-300 psig. These conditions are obtained
within specially designed heating coils. The heating coils
183
-------
-e- GAS
oo
/\
HF.DUCCD
CRUOF CHARGE
GASOLINE
STEAM
GAS OIL
7D-15?3-I
Figure F4-6. Process flow diagram - visbreaking.
Source: Reference 2
-------
can be arranged to provide a soaking section of ^relatively low
heat density. The charge remains in the coils until the vis-
breaking reactions are completed.1
An example of typical charge and products specifica-
tions is given in Table F4-14. Operating information and utility
requirements are included in Table F4-15.
TABLE F4-14. TYPICAL VISBREAKING CHARGE AND PRODUCT PROPERTIES
Charge
Gravity, API
Pour Point, °F
Viscosity, CS at 122CF
CS at 210°F
Products, Wt %
Gas
Unstabilized Gasoline (350°F E.P.)
350-650°F Gas Oil
650°F+ Residue
Properties of 350°F+ Visbreaker Product
"API
Viscosity, CS at 210°F
Pour Point, °F
Atmospheric
Residue
17.7
50.0
175.0
22.0
2.5
7.5
24.0
66.0
21.5
10.0
40. 0
Vacuum
Residue
6.6
120.0
1900.0
-
2.5
7.5
14.0
76.0
10.6
380.0
105.0
Source: Reference 4
185
-------
TABL-ErF4-15. VISBREAKER OPERATING AND UTILITY INFORMATION
Temperature, °F 850-950
Pressure, psig _ 50-300
Utilities - per barrel feed
Electricity, kWh 0.47
Steam (300 psig), Ib 8.7
Fuel, 103 Btu 88
Source: References 1, 4
4.2.3.2 Process Emissions--
Emission sources from visbreaking units include:
• Process heater flue gas, and
• Fugitive emissions.
Process Heater Flue Gas — Emission factors for various
pollutants from fired heaters are given in Table F4-16. A
typical visbreaking unit will utilize one process heater. It is
used to heat the feed to temperatures required for the visbreak-
ing reactions to proceed.
Total emissions from the visbreaker heater are also
given in Table F4-16. These figures are given as pounds of
pollutant per thousand barrels of fresh feed using a heat input
of 9.0 x lO1* Btu per barrel of feed.4
186
-------
TABLE F4-16.
TYPICAL EMISSIONS FROM VISBREAKING
UNIT PROCESS HEATERS -
EPA Emission Factor
(lb/103 gal-oil fired)
(lb/106 scf-gas fired)
Total Emissions
(lb/103 bbl of
fresh feed)
Oil Fired Heaters
Particulates
- Distillate oil
- Residual oil
Grade 4
Grade 5
Grade 6
Sulfur Dioxide0
- Distillate oil
- Residual oil
Sulfur Trioxidec
Carbon Monoxide
Hydrocarbons (as
Nitrogen Oxides
(as N02)
- Distillate oil
- Residual oile
7
10
10(S)+3
142(S)
157(S)
2(S)
5
1
22
22+40000
1.3
4.5
6.4
6.4(S)+1.9
91.3(S)
101(5)
3.2
0.64
14
14+257(N)'
Gas Fired Heaters
Particulates '
Sulfur Oxides (as S02)f
Carbon Monoxide
Hydrocarbons (as CHu)
Nitrogen Oxides (as N02)
5-15
0.6
17
3
120-230
0.43-1.29
0.051
1.5
0.26
10.3-19.7
Source: Reference 15
Based on a heat input of 9.0x10'' Btu/bbl of fresh feed with the following
fuel heating values: Oil - 140,000 Btu/gal; Gas - 1050 Btu/scf.
CS - Ht X sulfur in the oil
Improper combustion may cause a significant increase in emissions
'use this emission factor for residual oils with less Chan 0.5Z CN<.5) nitro-
gen content. For oil with higher nttrogen content Qj>0,5i, use emission
factor of 120 lb/103 gal
Based on sulfur content of 2000 gr/105 scf
187
-------
7 Fugitive Emissions — Fugitive emissions can occur from
a variety of sources including valves, pumps, compressors,
flanges and other fittings, relief valves, and drains. Hydro-
carbon emission rates for each individual source are quite low.
However, total fugitive emissions are significant because of the
large number of such sources within the refinery.
Total estimated fugitive non-methane hydrocarbon emis-
sions from a typical visbreaking unit are given in Table F4-17.
The listed emission factors were determined as a part of this
program. Additional information on these emission factors, in-
cluding a complete discussion on their derivation and the con-
fidence intervals for each source category, are contained else-
where in this report (see Appendices B and C).
Estimates for the number of sources within each source
category were developed from:
• Actual field source counts conducted during
the course of this program, and
• Counts contained in The Development of Petroleum
Refinery Plot Plans, EPA publication number
450/3-78-025.u
The methods used to develop estimates for the number
of each source type have been discussed in Section 4.1.1.3 of
this appendix.
The source counts described in Table F4-17 refer to
those sources located within the battery limits of the process.
That is, equipment located in tankage or transfer lines is not
include^ in these source counts.
186
-------
TABLE F4-17.
ESTIMATED FUGITIVE NON-METHANE HYDROCARBON EMISSIONS
FROM A TYPICAL VISBREAKING UNIT
\ 1
Emissions
Source Type
Valves
Open- End
(Sample)
Valves
Pumps (Pump
,_, Seals)
00
vD
Drains
Flanges 6
Fittings
Relief Valves
Compressors
(Compressor
Seals)
Process
Stream Service
Classification
Gas /Vapor
Light Liquid
(v? > o.i p«i« e iflo'F)
Heavy Liquid
(VP < 0.1 pai« 9 lOO'F)
Hydrogen Service
Total
All
Light Liquid
(V? > 0.1 paU 9 100'F)
Heavy Liquid
(VP < 0.1 p«i« 8 100°P>
Total
All
All
All
Hydrocarbon
Hydrogen
Total
Number of Sources in Process Unit
30
46
224
0
300a
20a
2t 2)
7(10)
9(12)a
23a
1071a
6a
0
0
0
Source
Emission
Factor, Ib/hr
0.059
0.024
0.0005
0.01B
0.005
0.25
0.046
0.070
0.00056
0.19
1.4
0.11
Estimated Total
Emissions,
Ib/hr
1.77
1.10
0.112
0.0
2.98
0.100
0.50
0.46
0.96
1.61
0.60
1.14
0.0
0.0
,0.0
7.29
Estimated
-------
- ~- The number of sources per unit depends on a variety
of factors including processing complexity, process type, and
processing flexibility. Hence, these source counts are not
necessarily representative of all or even the majority of vis-
breaking units.
4.2.4 References for Section 4.2
1. Hobson, G.D., and W. Pohl. Modern Petroleum Technology.
Applied Science Publishers Ltd., Essex, Great Britain,
1975.
2. Gary, James H. and Glenn E. Handwerk. Petroleum Refining,
Technology and Economics. Dekker, New York, 1975.
3. Cantrell, Ailleen. Annual Refining Survey. Oil and Gas J.,
77(13): 122-156, March 26, 1979.
4. Hydrocarbon Processing Refining Process Handbook. 57(9):
103, September 1978.
5. Rose, K. E. Delayed Coking--What You Should Know. Hydro.
Proc., 50(7): 85-92, July 1971.
6. Reis, Thomas. To Coke, Desulfurize and Calcine. Hydro.
Proc., 54(6), June 1975.
7. Nelson, W. L. Petroleum Refinery Engineering. McGraw-Hill
Book Company, Inc., New York, 1958.
8. Busch, Robert G. Fluid Coking. Oil and Gas J., 68(14):
102-111, April 6, 1970.
9. Flexicoking: An Advanced Fluid Coking Process. API Proc.
Ref., 1972.
10. Foulkes, P. B. and M. D. Harper. Prospects for Coking May
Brighten. Oil and Gas J., 76(12): 85-93, March 20, 1978.
11. Radian Corporation. Control Techniques for Carbon Monoxide
Emissions. Final Report. EPA-450/3-79-006, Austin, Texas,
June 1979.
12. Flexicoking Passes Major Test. Oil and Gas J., 73(10): 53-
56, March 10, 1975.
190
-------
13. National Air Pollution Control Techniques-Advisory Com-
mittee. Control Techniques for Carbon Monoxide Emissions
from Stationary Sources. NAPCA Pub. No. AP-65, March 1970.
14. Environmental Protection Agency. Development of Petroleum
Plot Plans. EPA-450/3-78-025, Research Triangle Park,
North Carolina, 1978.
15. Environmental Protection Agency. Compilation of Air Pol-
lutant Emission Factors. Third Edition, Supplement No. 8.
AP-42. Research Triangle Park, N.C., May 1978.
16. Air Pollution Control District, County of Los Angeles;
State of California, Department of Public Health; U.S.
Department of Health, Education and Welfare; Western Oil
and Gas Association. Emissions to the Atmosphere from
Eight Miscellaneous Sources in Oil Refineries. June 1958.
191
-------
4.3 CRACKING OPERATIONS " -
The primary function of cracking operations is to
convert heavy petroleum fractions into lighter,--more valuable
products. Two processes, catalytic cracking and hydrocracking,
provide a substantial portion of the cracking capacity in the
United States. Although these processes are similar in that
they crack heavy fractions to produce lighter products, there
are considerable differences between them in both the operating
principles and the pollution potential.
Catalytic cracking is a high temperature-low pressure
process which produces a relatively unsaturated product. The
removal of contaminants such as sulfur and nitrogen compounds
from the product streams is low and the conversion of feed to
lighter products is generally limited by the level of heavy
aromatics in the feed. If contaminant removal or higher con-
versions are required, hydrogenation of feed, product, or
recycle streams may be incorporated into the processing scheme.
More detailed information on catalytic cracking is provided in
Section 4.3.1.
Hydrocracking is a high temperature-high pressure
process which consumes hydrogen to produce highly saturated,
contaminant free products. Sulfur and nitrogen compounds in
the feed are converted to hydrogen sulfide and ammonia.
The hydrocracking process is characterized by its
flexibility. By varying the severity of the process conditions,
a variety of feedstocks, including those with high aromatics
concentrations, can be processed to yield a wide range of
product streams. Detailed information on hydrocracking pro-
cesses ii given in Section 4.3.2.
192
-------
The decision by the refiners as to which of these
processes to employ is complicated by a number o'f factors.
These include the availability of hydrogen, the desired product
slate, product quality requirements, and the degree of flexibil-
ity desired. Since each process has its own particular require-
ments for downstream processing, the choice of one process over
the other is usually based on an economic evaluation covering
the operation of the entire refinery.
A.3.1 Catalytic Cracking
The catalytic cracking process often serves as the
heart of the modern refinery by substantially increasing the
production of gasoline from a given amount of crude oil. This
is accomplished by cracking heavy feedstocks such as atmospheric
or vacuum gas oils to produce slurry oil, light cycle oil,
cracked gasoline, light gases, and coke.
Several types of catalytic cracking processes have
been developed: fluid-bed catalytic cracking (FCC) units, and
moving bed designs such as Thermofor (TCC) and Houdriflow (HCC)
cracking units. Table F4-18 gives a breakdown by process type
of catalytic cracking capacity in the United States as of
January 1979.
TABLE F4-18. DOMESTIC CATALYTIC CRACKING CAPACITY, 1979
Fresh Feed % of Total # of Units
Type Bbl/Sd Feed Capacity in Operation
FCC 4,592,500 94.0 119a
TCC 242,900 5.0 19
HCC 51,500 1.0 3
aThis value is somewhat low as some refineries have multiple
catalytic cracking units.
Source:- Reference 2
193
-------
4.3.1.1 ^-Process Description--
Figure F4-7 shows a diagram of a typical FCC unit.
Hot regenerated catalyst, mixed with hydrocarbon feed, is
transported into the reactor. The reactor contains a bed of
powdered catalyst which is kept in a fluidized state by the
flow of vaporized feed material and steam. Cracking of the
feed, which occurs in both the riser leading to the reactor
and in the fluidized bed, causes a deposit of coke to form on
the catalyst particles. A continuous stream of spent catalyst
is withdrawn from the reactor. The catalyst is steam stripped
to remove hydrocarbons and conveyed to the regenerator by air-
flow. The hydrocarbon vapor from the reactor is fractionated
into a variety of products including light hydrocarbons, cracked
gasoline, and fuel oil, while a portion of the fractionator
bottoms can be recycled to the reactor.3
Air is injected into the regenerator to burn off the
coke deposit and the regenerated catalyst is continuously
returned to the reactor. Heat added to the catalyst during
coke burn-off furnishes much of the required heat for the
cracking reaction.1*
Thermofor and Houdriflow catalytic cracking units
utilize beaded or pelleted catalysts. Regenerated catalyst
and vaporized feed enter the top of the reactor chamber and
travel cocurrently downward through the vessel. The catalyst
is purged with steam at the base of the reactor and travels by
gravity into the regenerator chamber. Combustion air is admitted
at a controlled rate to burn off coke deposits. From the bottom
of the regenerator, the catalyst is conveyed by airlift to a
surge hopper above the reactor. A diagram of a typical TCC unit
is given in Figure F4-8.
194
-------
vb
Cn
Electrostatic
Preclpltator
Srnum
Boiler Pcedvater
Regenerator
Catalyst
Fines
Mul t Ir.yclnnes
----
Gas
Oil
Gas
on
Figure F4-7. Fluid catalytic cracking unit (FCC).
Source: Reference 5
-------
VO
Gas Oil
Charge
*- Uet Gas
Catalytic
GasolIne
Light
Fuel
B- Heavy Funl
Figure F4-8. Thermofor moving-bed catalytic cracker (TCC).
Source: Reference 14
-------
- ~~ With the advent of new catalysts, major design and
operational changes have been incorporated in FCC unit opera-
tion. By contrast, no major changes in moving bed type units
have been observed and these units are being phased out.
4.3.1.2 Process Technology--
Conventional FCC Operation—Coke is removed from
cracking catalysts to restore their activity. Spent catalyst
contains roughly 6 percent coke while coke levels on the regen-
erated catalyst vary frorr. 0.2-0.3 percent. The amount of air
supplied to the regenerator is insufficient for complete com-
bustion which results in flue gas CO concentrations of 5-10
percent. The oxygen level in the flue gas is low enough so that
only a limited amount of combustion occurs in the regenerator
"dilute phase" where no catalyst heat sink is available. Com-
bustion in the dilute phase, called afterburning, can result in
damage to the catalyst, the cyclones, and other regenerator
equipment due to high temperatures. To avoid equipment damage,
the regenerator is operated beloxv 1150°F. 5 ' 6 ' 7
High Temperature Regeneration—Zeolite catalysts
first appeared on the market in the mid-19601s. The major
features of these catalysts are summarized below:6'8
• Naphthenes and paraffins are cracked
rapidly with excellent selectivity,
• Aromatic nuclei crack slowly with
poor selectivity,
• High hydrogen transfer rates are
observed,
197
-------
~~ • The rate of cracking is relatively
unaffected by boiling range, and
• Efficiency is adversely affected by
coke deposits which limit zeolite
availability.
The use of zeolite catalysts has accelerated the
trend to more fully regenerate these coke sensitive catalysts.
Very low carbon on regenerated catalyst (CRC) levels have been
achieved using a technique called high temperature regeneration
(HTR). The key to this process is complete conversion of coke
to C02 within the regenerator. This situation is quite differ-
ent from that of conventional regeneration where conversion of
CO to C02 is carefully controlled to about 50 percent.7 High
temperature regeneration can be utilized in new units, or
applied as a retrofit to existing units. The major features of
high temperature regeneration are:
• Extremely low coke on regenerated catalyst
levels are possible. Typical values are
0.05-0.1 percent coke while Amoco Oil
Company reported CRC levels of 0.01 percent
with their UltraCat regeneration technique.5'7
• CO emission levels of 500 ppm in the regen-
erator flue gas can be obtained. This level
is sufficiently low to meet federal NSPS and
most state emissions regulations.7'9':D':l
• Complete regeneration increases catalyst
activity which means a lower catalyst-to-oil
ratio is possible. Thus, unit capacity can
be increased if bottlenecks are removed from
the rest of the process.5'6'7'8'9
198
-------
- r" • Temperatures in the regenerator v-ary from
1100-1350°F. This is 100-150°F higher than
conventional regeneration. Since CO after-
burn is possible, flue gas temperatures can
be several hundred degrees higher than the
dense-phase temperature. Normally, however,
only 50 to 100'F of temperature differential
results.4>7>9'10
• The extremely active catalyst produced from
HTR is most effectively used in a short
contact time riser cracking reactor. The
advantage of riser cracking over bed cracking
lies in avoiding secondary reactions such as
the recracking of gasoline components. 6'9
• Increased catalyst selectivity and the use
of riser cracking can result in a 20 to 30
percent reduction in the amount of coke pro-
duced. Therefore, the increase in combustion
air required to completely burn CO and coke to
COa can be offset in some cases by lower coke
production -such that overall combustion air
usage can remain essentially constant.5
The operating conditions for conventional FCC's and
units using HRT are compared in Table F4-19.
Existing FCC units may be revamped to incorporate
KTR. The required changes to convert to HTR depend on the
design of the unit and the desired coke content on the regen-
erated catalyst. To withstand higher regenerator temperatures,
steel components within the regenerator usually require replace-
ment with components made of more heat resistant materials such
199
-------
as chromjtum-nickel alloy stainless steels. Other modifications
might include an improved combustion air distribution system
or the installation of a riser cracking reactor. In general,
switching to HTR increases the capacity of the process and some
modifications in downstream equipment may be required to remove
bottlenecks.8'9
TABLE F4-19. TYPICAL OPERATING CONDITIONS FOR FLUID
CATALYTIC CRACKING
Reactor Temperature, °F 885 - 1025
Regenerator Temperature, °F
Conventional Regeneration 1000 - 1100
HTR 1100 - 1350
Coke Content of Spent Catalyst, Wt %
Conventional Regeneration 6
HTR 5
Coke Content on Regenerated Catalyst, Wt %
Conventional Regeneration 0.2 - 0.3
HTR 0.01-0.1
Source.- References 5, 9
Combustion Promotion Catalysts--The most recent devel-
opment in FCC technology is the use of "promotion" catalysts to
aid the conversion of CO to COa. The first type to become
available (1975) was an FCC catalyst modified with a small con-
centration of noble metal promotion agent. In 1977, a number of
manufacturers came out with a solid promoter. This powder is
mixed with make-up catalyst, roughly 1-10 pounds per ton of
fresh catalyst. Liquid promoters, injected directly into the
regenerator, are also available.11
200
-------
? The advantage of using combustion promoters is that
CO is converted to €62 within the dense-phase of the regenerator.
This avoids the problem of CO afterburn in the regenerator dilute
phase which is common in units using HTR. Thus-, in units where
temperature limitations prohibit the use of HTR, CO emissions
below 500 ppm can be obtained without the use of a CO boiler.
Essentially complete conversion of CO can be obtained with bed
temperatures of 1150-1200°F.7'° However, regeneration of the
catalyst is not quite as effective at the lower temperature and
selectivity of the catalyst is slightly poorer in that more coke
is produced.9 Typically, FCC units with CO promoted regenera-
tion have regenerator bed temperatures around 1300°F and dilute
phase temperatures are generally 10 to 30°F higher.
Industry acceptance of high temperature regeneration
and promotion catalysts has been high, given the relatively
short time they have been available. A breakdown of the
regeneration techniques being used by U.S. refiners is given in
Table F4-20.
TABLE F4-20. DOMESTIC FCC REGENERATION TECHNIQUES (August 1978)
% of All FCC
Regeneration Units that Use
Technique This Technique Remarks
Conventional regeneration 53 Most units have CO boilers
High temperature regeneration 26 May be used in conjunction
with a CO boiler
Combustion promoting catalysts 10 May be used in conjunction
with a CO boiler
Combustion promotion, separate 11 May be used in conjunction
from catalyst with a CO boiler
Source: Reference 15
201
-------
- i-- Energy Factors--There are a number of-factors involved
in determining the energy requirements of the fluid catalytic
cracking process. The cracking reaction itself is e-ndothermic
and the heat of reaction is around 150 Btu/lb feed. The heat
of combustion of the coke provides the majority of the required
energy. That is, coke burning in the regenerator heats the
catalyst and this heat is transferred to the oil in the riser.
The remainder of the required heat is obtained by preheating
the feed in a fired heater.
The flue gas from conventional FCC operation contains
from 5-10 percent CO which represents a substantial energy loss
if released to the atmosphere. This energy can be recovered in
a CO boiler by producing steam. Energy recovery from a typical
CO boiler, operating with a conventional FCC unit, has been
estimated to be 60 x 103 Btu/bbl of fresh feed.12 Often the
entire cost of a CO boiler can be justified on the basis of
steam production alone.
Required utilities for the FCC process include fuel
for the fired heater and auxiliary fuel for the CO boiler (if
necessary). Steam is used in the stripper to strip oil from
the catalyst particles and additional energy is required to
power the regenerator air blower.
Energy recovery from high temperature regeneration is
about the same as for CO boilers, estimated at 60 x 103 Btu/bbl
of fresh feed.12 In high temperature regeneration, the combus-
tion of CO to COa occurs within the regenerator rather than in
the CO boiler. This energy produces higher flue gas tempera-
tures which results in more steam production in the waste heat
boiler, greater heat transfer co the catalyst which reduces the
preheate_r duty, and increased yields of more valuable liquid
products.
202
-------
i- Many refiners operate their FCC regenerator at higher
pressures (25 to 40 psig) to take advantage of higher CO com-
bustion rates resulting from the increased oxygen partial
pressure. Power recovery equipment, used to recover mechanical
energy from the higher pressure flue gas, has been installed on
some of these units. The system consists of a flue gas expander,
the FCC air blower, an induction motor-generator, and a start-up
steam turbine, all mounted in tandem. The flue gas passes through
an efficient third stage cyclone system to the flue gas expander
as shown in Figure F4-9. The expander produces power for the FCC
air blower. The induction motor-generator smoothes any imbalance
in power load by adding power if necessary or by producing elec-
tricity if the power generated by the flue gas expander is in
excess of air blower requirements. Efficiency and design improve-
ments in the third stage separator and the use of erosion-tempera-
ture resistant expander blades have made the system economically
attractive for new FCC units.5
The energy recovery potential of various power
recovery schemes are compared in Table F4-21.
4.3.1.3 Process Emissions--
Emission sources from catalytic cracking units
include:
• Catalyst regeneration operations,
• Process heater flue gas,
• Fugitive emissions.
203
-------
Regenerator
Cyclones -
Catalyst -
— Flue Gas
Atr
Orifice
Chamber
Motor A|r
Driver Dlower
To Stack
j
Boiler Feed
Water -
Regenerator
r
Catalyst
Fines
Steam
CO
flntler
ElectrostalIc
Preclpalator
To Stack
Electrostatic
1—| ' 1 ' t iccirosiau
L_Bollcr Feed! Preclpatator
Water t
Catalyst Fines
Air
Steam
Expander Motor/Gen
Start-up
, Steam
Turbine
^ 7-
Air to Regenerator _/
Air Olower Exhaust
Figure F4-9. FCC unit: power recovery system.
Source: Reference 13
-------
TABLE F4-21. ENERGY RECOVERY TECHNIQUES FOR FCC REGENERATORS
Available Energy-Horsepower per
Power Recovery Technique Pound per Second of Flue Gasa
Waste heat boiler alone 45
Power recovery alone 78
Power recovery plus waste heat 106
o boiler
Ul
CO burning waste heat boiler 145
Power recovery plus CO burning 206
waste heat boiler
aFlue gas at 20 psig - 1000°F.
Source: Reference 3
Note: This table assumes that the regenerator is operating in "conventional"
regeneration mode. If a complete combustion regeneration mode is used
(high temperature regeneration or CO promoted catalyst), then the
incremental recoveries shown for CO burning will be obtained in the
regenerator and in the unfired waste heat boiler.
-------
~ Catalyst Regeneration Operations--CokeLis deposited on
cracking catalysts during processing, and it must be removed to
restore catalytic activity and selectivity. This is accomplish-
ed by introducing air into the regenerator which burns the coke
to CO and C02.
In conventional operation, the conversion of CO to C02
is minimized to avoid high temperatures which might damage
internal regenerator materials. The resulting flue gas contains
from 5-10 percent CO.
Many refiners utilize a CO-burning waste heat boiler
to recover the energy contained in the flue gas. In addition to
energy recovery, this procedure reduces CO emissions to low
levels. And, other combustible contaminants may be removed,
depending on the operating temperature of the boiler. Increased
emissions of "thermal NOX" may, however, result at boiler
temperatures and sulfur in the auxiliary fuel will be converted
to SOx. Typical emission levels for regenerator flue gas with
and without the use of a CO boiler are listed in Table F4-22.
Also given in Table FA-22 are the results of sampling conducted
during this program. The table lists a summary of data
obtained from five FCC unit CO boiler stacks. Detailed informa-
tion on these sampling results is given in Appendix B of this
report.
Cracking units utilizing high temperature regeneration
produce a flue gas with CO concentrations ranging from 200-2000
ppm. And, many of the applicable governmental regulations have
been set or altered to allow the use of this technology.
206
-------
TABLE FA-22. EMISSION RATES FROM FCC REGENERATORS, BEFORE AND AFTER CO BOILER
Chcrolcnl Species
SOj, ppmv
SOj, ppmv
N0x (a.q N02), Ppmv
CO, % Vol.
C02, X Vol.
02, Z Vol.
N2, % Vol.
H20, °i Vol.
Hydrocarbons, pprav
Ammonia, ppmv
Aldehydes, ppmv
Cyanides, ppmv
Participates, gr/SCF
Temperature, °F
limloslona without
CO Bnilpr
(Keference 9)
130-3300
NA'
8-394
7.2-12.0
10.5-11.3
0.2-2. /i
78.5-80.3
13.9-26.3
98-1213
0-675
3-130
0.19-0.94
O.Oli-1.39
1000-1200
(Reference 9)
Up to 2700
NAC
Up to 500
0-14 ppmv
11.2-14.0
2.0-6.4
82.0-84.2
13.4-23.9
NAC
NAC
NAC
NAC
0.017-1.03
485-820
('.missions with
CO Boiler
D.ita from , Total
Current Prograw Current
14.4-871
0.65-13.5
94.1-453
0.0-1.0
13.5-16.1
3.2-7.0
77-82.7
9.2-22.7
0.28-46.2
0.0-15.4
0.0-19.6
0.0-19.1
0.012-0.304
386-727
Emlstjlona Baucd on Data from
Program - (lb/1000 bbl feed)
9-11-504
0.37-10.3
37.2-215
0.0-2680
-
-
-
-
0.17-29.8
0.0-2.37
0.003-5.31
0.0-4.66
6.8-236
-
.All concentrations on dry bnala
Baaed on sampling of 5 stacks
GNot available
-------
High temperature regeneration can reduce the uncon-
trolled emissions of other pollutants. Since the maximum
temperatures used in high temperature regeneration are well
below those used in CO boilers, thermal NOX emissions are
somewhat lower.9 And, nitrogen and sulfur emissions from CO
boiler auxiliary fuel use are avoided.
CO emissions from various operating schemes are
summarized in Table F4-23.
TABLE FA-2 3.
SUMMARY OF CO EMISSIONS FROM VARIOUS
REGENERATION TECHNIQUES
Control Technique
Typical Emissions Level
Conventional regeneration,
(uncontrolled) :
Conventional regeneration,
(CO boiler)
High temperature regeneration
or combustion promotion
5-10% CO in regenerator flue gas; AP-42
emission factor=13,700 Ib CO/1000 bbl feed
<50 ppm CO in CO boiler flue gas
200-2000 ppm CO in regenerator flue gas;
<500 ppm CO can usually be obtained
Source: Reference 9
Process Heaters — Emission factors for various pollu-
tants from fired heaters are given in Table F4-24. A typical
catalytic cracking unit will require a process heater for heat-
ing the feed, prior to its contact with catalyst in the riser
and reactor.
203
-------
TABLE F4-24.
TYPICAL EMISSIONS FROM CATALYTIC CRACKING
UNIT PROCESS HEATERS
EPA Emission Factor
(lb/103 gal-oil fired)
(Ib/lO* scf-gas fired)
Total Emissions
(lb/103 bbl of
- fresh feed)
Oil Fired Heaters
Particulates
- Distillate oil
- Residual oil
Grade 4
Grade 5
Grade 6
Sulfur Dioxide
- Distillate oil
- Residual oil
Sulfur Trioxidec
Carbon Monoxide
Hydrocarbons (as CH^)'
Nitrogen Oxides
(as N02)
7
10
10(S)+3
142(S)
157(S)
2(S)
5
1
1.4
5.0
7.1
3.6
0.7
- Distillate oil
- Residual oil6
Gas Tired Heaters
Particulates
Sulfur Oxides (as S02>
Carbon Monoxide
Hydrocarbons (as CE\)
Nitrogen Oxides (as N02)
22
22+400 (N)2
5-15
0.6
17
3
120-230
15.7
15. 7+286 (N)2
0.5-1.4
0.06
1.6
0.29 '
11.4-21.9
Source: Reference 17
Based on a heat input of 100,000 Btu/bbl of fresh feed with the following
fuel heating values: Oil - 140,000 Btu/gal; Gas - 1050 Btu/scf.
CS - Wt Z sulfur in the oil
Tmproper combustion may cause a significant increase in emissions
^se this emission factor for residual oils with less than 0.5Z (N0,5), use emission
factor of 120 Ib/lO* gal
Based on sulfur content of 2000 gr/106 scf
209
-------
- r Total emissions from the catalytic crocking unit
heaters are also given in Table F4-24. These figures are given
as pounds of pollutant per thousand barrels of fresh feed based
on a representative heat input of 100,000 Btu per barrel of
fresh feed.15)28 -
Fugitive Emissions — Fugitive emissions can occur from
a variety of sources including valves, pumps, compressors,
flanges and other fittings, relief valves, and drains. Hydro-
carbon emission rates for each individual source are quite low.
However, total fugitive emissions are significant because of the
large number of such sources within the refinery.
Total estimated fugitive non-methane hydrocarbon emis-
sions from a typical catalytic cracking unit are given in Table
F4-25. The listed emission factors were determined as a part of
this program. Additional information on these emission factors,
including a complete discussion on their derivation and the con-
fidence intervals for each source category, are contained in
Appendices B and C of this report.
Estimates for the number of sources within each source
category were developed from:
• Actual field source counts conducted during the
course of this program, and
• Counts contained in The Development of Petroleum
Refinery Plot Plans, EPA publication number 450/
3-78-025.18
The methods used to develop estimates r£or the number
of each^source type have been described in Section 4.1.1.3 of
this appendix.
210
-------
TABLE F4-25.
ESTIMATED FUGITIVE NON-METHANE HYDROCARBON EMISSIONS
FROM A TYPICAL CATALYTIC CRACKING UNIT
Emissions
Source Type
Valves
Open-End
(Sample)
Valves
Pumps (Pump
Seals)
Drains
Flanges &
Fittings
Process
Stream Service
Classification
Gas/Vapor
Light Liquid
(VP > 0.1 psla 9 100'P)
Heavy Liquid
(VP < 0.1 psla 9 100'P)
Hydrogen Service
Total
All
Light Liquid
(VP > 0.1 psla ? 100'P)
Heavy Liquid
(VP < 0.1 psla 9 100'F)
Total
All
All
Relief Valves All
Compressors
(Compressor
Seals)
Hydrocarbon
Hydrogen
Total
Number of Sources
In Process Unit
Counts or Estimates Counts or Estimates Emission
From Radian Study From PES Study Factor, Ib/h
384
409
521
0
1314a
_
13(18)
17(24)
30(42)a
65a
4214*
6C
4(8)
0
4(8)a
849
889
1167
0
2905C
67b
16(22)
21(29)
37(52)b
-
9635C
-
4(8)
0
4(8)t>
0.059
0.024
0.0005
0.018
0.005
0.25
0.046
0.070
0.00056
0.19
1.4
0.11
Estimated Total,
Emissions.
r Ib/hr
22.7 - 50.1
9.82 - 21.3
0.261 - 0.584
0.0
32.8 - 72.0
0.335
4.5 - 5.50
1.10 - 1.33
5.60 - 6.83
4.55
2.36 - 5.40
1.14
11.2
0.0
58.0 - 101
Physically Counted
Counted From Flow Diagrams
Estimated
Reference 18
-------
~ The source counts given in Table F4'-25Lrefer to those
sources located within the battery limits of the process. That
is, equipment located in tankage or transfer lines is not
included in these source counts.
The number of sources per unit depends on a variety
of factors including processing complexity, process types, and
processing flexibility. The source counts presented in Table
F4-25 are not necessarily representative of all or even the
majority of catalytic cracking units.
The estimated composition of fugitive non-methane emis-
sions from a catalytic cracking unit is given in Table F4-26.
These estimates indicate the types of hydrocarbons contributed by
sources on a number of process streams found within the unit.
Additional information about these estimates may be found in
Section 2.4 of Appendix D.
4.3.2 Hydrocracking
Hydrocracking refers to the process of converting
heavy feedstocks into lighter products in the presence of hydro-
gen and a suitable catalyst or series of catalysts. Hydro-
cracking applications extend from cracking naphtha for LPG
production to cracking heavy gas oil or residuum for products
including gasoline, catalytic cracking feedstock, and desulfur-
ized fuel oils.21* The products of the hydrocracking process are
highly saturated, that is, non-olefinic, and are characterized
by low levels of sulfur and nitrogen. Generally, branched
paraffin isomers predominate and very little methane or ethane
is formed.2S
212
-------
TABLE F4-26.
N)
M
CO
ESTIMATED COMPOSITION OF FUGITIVE NON-METHANE HYDROCARBON
EMISSIONS FROM A FLUID CATALYTIC CRACKING UNIT
estimated percentage of emissions
attributed to each stream - 1.
Weighted contribution of each
component to unit emissions - ppraw
Benzene
Toluene
Ethylbenzene
Xylcncs
Other Alkylbenzenea
Naphthalene
Anthracene
Biphenyl
Other Polynuclear aroma tics
n-Hexane
Other Alkanes
Olefins
Cyclo Alkanea
Hydrogen
Atmospheric
GAS Oil
1
0
0
0
0
1
0
0
0
2
0
9495
0
500
0
Stream
Fuel
Gas
30
0
0
0
0
0
0
0
0
0
0
216000
18000
0
6000
LPG
Olefins
23
0
0
0
0
0
0
0
0
0
0
92000
138000
0
0
Cracked
Naphtha
45
1296
40401
9644
77153
109562
4928
0
0
2916
5324
91850
76833
30096
0
Lt. Cycle
Gas Oil
1
0
0
0
6
267
590
103
102
6245
0
1906
368
412
0
Hvy. Cycle
Gas Oil
0
0
0
0
0
0
0
0
0
0
0
0
0
0
1 0'
Totals
100Z
1296
40401
9644
77159
109830
5518
103
102
9163
5324
471251
233201
31008
6006
1000000
-------
... Catalytic reforming of the gasoline boiling range
product'is usually required to produce high octane gasoline.
However, the hydrocracking product makes excellent reformer
feedstock and the hydrogen produced in the reformer can provide
much of the hydrogen required in the hydrocracker.
The total hydrocracking capacity in the United States
has shown only limited growth in the last few years. This is
due primarily to the development of new catalysts for the fluid-
bed catalytic cracking process, and the high hydrogen require-
ments of hydrocracking units.22 Today, hydrocracking is being
used as a companion process to catalytic cracking. That is,
feedstocks which would present difficulties for catalytic
cracking are processed in the hydrocracker. Examples of such
feedstocks include coker gas oils and catalytic cracker cycle
oils which are high in aromatics.2°'26 Current hydrocracking
capacity in the United States is given in Table F4-27.
TABLE F4-27. U. S. HYDROCRACKING CAPACITY, 1979
Feedstock
Distillate Streams
Residual Streams
Other
Number of
Units
37
1
4
42
Total Capacity,
bbl/sd
755,400
20,000
110,000
885,400
Source: Reference 2
Several types of reactions occur during hydrocracking.
The first group of reactions results in the removal of sulfur
and nitrogen impurities as indicated below:23
214
-------
i- RS + H2 * R + R2S ....^
2RN + 3H2 + 2R + 2NH3
This is followed by the cracking and hydrogenation
reaction:2 3
R + H2 r* R1 + C, + C2 + etc
Hydrocracking catalysts serve two functions. They
are (1) cracking of high molecular weight hydrocarbons, and
(2) hydrogenation of the unsaturates formed during the cracking
step or otherwise present in the feedstock.
A typical hydrocracking catalyst contains a cracking
component such as silica-alumina. The silica alumina may be in
the form of acid treated aluminosilicates, amorphous synthetic
silica-alumina combinations, or the silica-alumina combinations
known as zeolites or molecular sieves. These materials are all
high temperature acids and their cracking activity is attributed
to their acidity.
The hydrogenation activity of hydrocracking catalysts
is obtained by incorporating a base metal or noble metal on the
silica-alumina base. Typical hydrogenating metals include nickel,
palladium, molybdenum, and cobalt.23 These may be altered by
promotion with another metal or by pretreatment techniques such
as sulfiding.20 >21
4.3.2.1 Process Description
Hydrocracking processes can be classified according to
differences in the process design and the type erf catalyst employed,
215
-------
Two different types of hydrocracking processes..are evident.
In one, the feedstock is severely pretreated over a hydrogenat-
ing catalyst of low acidity to remove essentially all of the
sulfur and nitrogen from the feed. Hydrocracking is carried
out after ammonia removal in a separate stage over a relatively
acidic cracking catalyst. This type of processing is particu-
larly useful with high nitrogen containing feedstocks since
nitrogen, present in both the feed and the resulting ammonia,
tends to poison cracking catalysts.
In the other approach, zeolite catalysts are employed
which can accomplish substantial cracking in the presence of sul-
fur and nitrogen compounds.27 This allows single stage operation.
Although there is a drop in catalytic activity over that of the
two stage design, less equipment is required and recompression
costs can be avoided.21
The process flow diagram for a typical single stage
hydrocracker is shown in Figure F4-10. Heavy oil feed is mixed
with recycle and make-up hydrogen. The mixture is heated and
charged to the reactor where substantial cracking and hydrogen-
ation of the feed occurs. In addition, sulfur and nitrogen
compounds in the feed are converted to hydrogen sulfide and
ammonia. The product mixture is sent to a high pressure sepa-
rator where hydrogen is recovered and recycled to the reactor.
The liquid stream is fractionated into a number of products
and a heavy uncracked stream is recycled to the reactor. In
some units, 100 percent conversion from feed to lighter products
is obtained.
A flow diagram for the two stage process is given
in Figure F4-11. In the two stage design, the effluent from
the first reactor is flashed, allowing separation of liquid
216
-------
11
SINGLE-STAGE REACTORS
/\
FURNACE
FRESH FEED,
ri.AS/l
GAS
HYDROGEN
RECYCLE
_TO C,. C5-C
RECOVERY
. C? NAPHTHA
^ JET OR
DIESEL FUEL
FRACT10NATOR
ESSORf
A,
MAKE-UP
HYDROGEN
COMPRESSOR
LIQUID RECYCLE OR
ADDITIONAL PRODUCT
70-1514-1
Figure F4-10. Process flow diagram - single stage hydrocracker,
Source: Reference 27
-------
FIRST-STAGE REALTORS
SECOND-STAGE REACTOR
FRACTIONATOR
No
M
00
FRFSII FFED ^ »
TO C,, C5-C6
RECOVER*
C? NAPHTHA
JET OR
DIESEL FUEL
LIQUID RECYCLE OR
ADDITIONAL PRODUCT
70-1515-1
Figure F4-11. Process flow diagram - two stage hydrocracker.
Source: Reference 27
-------
from light gases containing the hydrogen sulfide. and ammonia
formed in the first stage. The lack of nitrogen""compounds in
the second stage feed allows the reactor to obtain a.high degree
of cracking at conditions of lower severity.
Several licensed hydrocracker processes have been
offered to the industry, many of which are available in both
one and two stage designs. Some of the more prevalent designs
are listed in Table F4-28.
TABLE F4-28. LICENSED HYDROCRACKING PROCESSES
Process Licensor
Isocracking Chevron Research Co.
Unicracking-JHC Union Oil Cc. and Esso Research
H-Oil Hydrocarbon Research Inc. and Cities
Service Research and Development Co.
H-G Hydrocracking Gulf Research and Development Co. and
the Houdry Div. of Air Products and
Chemicals Inc.
Hycracking Exxon Research and Development Co.
Most of these processes use a fixed bed reactor where
the feed enters the top of the unit and travels downward through
the catalyst bed. An exception to this is the H-Oil process
which uses an ebullient bed reactor. In the H-Oil process,
catalyst is continuously withdrawn and regenerated outside of
the reactor. Thus, feeds which contain high concentrations of
metals, sulfur, or nitrogen can be processed. The H-Oil
process is commonly used to upgrade residuum. l>t
219
-------
4.3.2,2;,- Process Technology--
The severity of the process conditions required for
hydrocracking depends on the type of feedstock and the degree
of cracking desired. In general, operating conditions must be
more severe when processing a heavy feedstock or when cracking
to increasingly lighter products.
Process Variables—The primary reaction variables
are the reactor temperature and pressure, and the nitrogen and
sulfur concentrations in the feed and off gases. The effects
of these variables are discussed below.3
Reactor temperature is the primary means of control
over the level of conversion. At normal reactor conditions, a
20°F increase in temperature nearly doubles the reaction rate.
However, the effect on conversion is somewhat less as substan-
tial recracking of material which has already been converted
can occur. Since catalyst activity decreases with time, it is
usually necessary to increase the temperature by 0.1 to 0.2°F
per day, depending on the rate of catalyst deactivation, over
the duration of the cycle.
An increase in reactor pressure serves to increase
the partial pressures of hydrogen and ammonia. Conversion
increases with increasing hydrogen partial pressure and de-
creases with increasing ammonia partial pressure. The hydrogen
effect is greater, however, and the net effect of an increase
in total pressure is an increase in conversion.
The organic nitrogen content of the feed is important
since hydrocracking catalysts are deactivated by organic
nitrogen compounds. A decrease in catalytic activity decreases
conversion.
220
-------
~ Low concentrations of hydrogen sulfide-_can serve as
an inhibitor for reactions involving the saturation of aromatic
rings.3 This reduces hydrogen consumption and prevents the
destruction of high octane aromatics. Higher hydrogen sulfide
concentrations, however, can decrease cracking activity and may
result in increased equipment corrosion rates.
Typical operating conditions and information on pro-
cess utility requirements are given in Table F4-29.
TABLE F4-29. HYDROCRACKING - TYPICAL PROCESSING CONDITIONS
AND UTILITY DATA
Pressure, psig 1000 - 4000
Temperature, °F 400 - 850
Hydrogen recycle, scf/bbl feed 8000 - 15000
Hydrogen consumption, scf/bbl feed 1500 - 2500
Fuel, 103 Btu/bbl feed 100 - 250
Power, kwh/bbl feed 6-15
Space velocity, v/hr/v 0.2 - 1.0
Source: References 3, 14, 20, 24
The high pressures required for hydrocracking result
in special problems when designing the reactors. In addition
to high pressures, several other factors add to reactor design
difficulties. These include.-
• High hydrogen partial pressures,
• High H2S and NH3 partial pressures,
221
-------
^- • High operating and regenerating
temperatures, and
• High exothermic heat of reaction.
The heat of reaction depends on the amount of hydrogen
absorbed and usually ranges between 55-70 Btu/scf H2. For a feed
absorbing 2000 SCF H2/bbl, the heat of reaction would be about
AGO Btu/lb feed. This heat can be removed from the reactor by
one of several methods : * *
• Injection of cold recycle gas,
• Injection of cold recycle liquid, or
• Inter-bed exchangers.
The first method, which cools the reaction mixture by
increasing vaporization, is most favored. Often, several reac-
tors are used for each stage to facilitate injecting the recycle
gas.3 Liquid recycle causes difficulty in proper liquid distribu-
tion across the bed, and inter-bed exchangers raise difficult
mechanical design problems under the high severity conditions
in the reactor.
4.3.2.3 Process Emissions--
Emissions from hydrocracking processes include:
• Emissions during periodic catalyst
regeneration,
222
-------
- • Process heater flue gas, and .
• Fugitive emissions.
Emissions from Catalyst Regeneration Operations--
During the hydrocracking process, deposits form on the catalyst
which must be removed to restore catalyst activity and selec-
tivity. This regeneration procedure involves burning the
deposits by admitting air into the reactors.
During regeneration, large quantities of carbon
monoxide and other pollutants may be released. However,
regeneration may be required only after several months or even
years of operation. Hence, total averaged emissions from this
source are generally insignificant.
Process Heater Flue Gas—Emission factors for various
pollutants from fired heaters are given in Table F4-30. A
typical hydrocracker will require a process heater to heat the
combined hydrocarbon-hydrogen feed to the reactor. In addition,
a process heater may be required as a reboiler for the main
fractionator. Heat input requirements vary, ranging from
100,000 to 250,000 Btu per barrel of oil feed.1"'20
Total estimated emissions from typical hydrocracker
heaters are also given in Table F4-30. These figures are given
as pounds of pollutant per thousand barrels of hydrocracker
feed using a total required heat input of 200,000 Btu per
barrel of feed.
Fugitive Emissions--Fugitive emissions can occur from
a variety of sources including valves, pumps, compressors,
flanges and other fittings, relief valves, and drains. Hydro-
carbon emission rates for each individual source are quite low.
223
-------
. TABLE F4-30.
TYPICAL EMISSIONS FROM HYDROCRACKING
UNIT PROCESS HEATERS ~.
EPA Emission ?actora
(lb/103 gal-oil fired)
(lb/106 scf-gas fired)
Total Emissions
(lb/103 bbl of
Jresh feed)
Oil Fired Heaters
Farticulates
- Distillate oil
. - Residual oil
Grade 4
Grade 5
Grade 6
Sulfur Dioxide
- Distillate oil
- Residual oil
Sulfur Trioxidec
Carbon Monoxide
Hydrocarbons (as CHj,)'
Nitrogen Oxides
(as N02)
- Distillate oil
- Residual oil6
7
10
10(S)+3
142(S)
157(S)
2(S)
5
1
22
22-KOO(N)'
2.9
10
14
14(S)+4.3
203(S)
224(S)
2.9(S)
7.1
1.4
31
31+571(N)"
Gas Fired Heaters
Participates
Sulfur Oxides (as S02)
Carbon Monoxide
Hydrocarbons (as CH^)
Nitrogen Oxides (as N02)
5-15
0.6
17
3
120-230
0.95-2.
0.11
3.2
0.57
22.9-43.
9
8
aSource: Reference 17
Based on a beat input of 200,000 Btu/bbl of fresh feed with the following
fuel heating values: Oil - 140,000 Btu/gal; Gas - 1050 Btu/scf.
CS - Wt Z sulfur in the oil
Improper combustion may cause a significant increase in emissions
llse this emission factor for residual oils with less than 0.5Z (N^.S} nitro-
gen content. For oil with higher nitrogen content 0^0,5), use emission
factor of 120 lb/103 gal
Based on sulfur content of 2000 gr/106 s-cf
224
-------
However*-" total fugitive emissions are significant because of the
large number of such sources within the refinery.
Total estimated fugitive non-methane, hydrocarbon emis-
sions from a typical hydrocracking unit are given in Table F4-31.
The listed emission factors were determined as a part of this
program. Additional information on these emission factors, in-
cluding a complete discussion on their derivation and the con-
fidence intervals for each source category, are contained in
Appendices B and C of this report.
Estimates for the number of sources within each source
category were developed from:
• Actual field source counts conducted during
the course of this program, and
• Counts contained in The Development of Petroleum
Refinery Plot.Plans, EPA publication number
450/3-78-025.18
The methods used to develop estimates for the number
of each source type have been previously described in Section
4.1.1.3 of this appendix.
The source counts listed in Table F4-31 refer to those
sources located within the battery limits of the process. That
is, equipment located in tankage or transfer lines is not
included in these source counts.
The number of sources per unit depends on a variety of
factors including processing complexity, process type, and
processing flexibility. The source counts given in Table F4-31
are not necessarily representative of all or even the majority
of hydrocracking units.
225
-------
TABLE F4-31.
N>
N>
ESTIMATED FUGITIVE NON-METHANE HYDROCARBON EMISSIONS FROM
A TYPICAL HYDROCRACKING UNIT
Emissions
Source Type
Valves
Open-End
(Sample)
Valves
Puropa (Pump
Seals)
Drains
Flanges &
Fittings
Relief Valves
Compressors
(Compressor
Seals)
Process
Stream Service
Classification
Gas/Vapor
Light Liquid
(VP > 0.1 pala 3 100'F)
Heavy Liquid
(VP < 0.1 pala 8 100'P)
Hydrogen Service
Total
All
Light Liquid
(VP > 0.1 pa la 9 100'F)
Heavy Liquid
(VP < 0.1 pala f 100'F)
Total
All
All
All
Hydrocarbon
Hydrogen
Total
Number of Sourceu in Process Unit
175
375
307
75
931a
-
12(17)
10(14)
22(31)a
58a
1955b
6b
0(0)
3(6)
3(6)a
Source
Emission
Factor, Ib/hr
0.059
0.024
0.0005
0.018
0.005
0.25
0.046
0.070
0.00056
0.19
1.4
0.11
Estimated Total
Emissions,
Ib/hr
10.3
9.00
0.154
1.35
20.8
-
4.25
0.644
4.9
4.1
1.09
1.1
0.0
0.66
, 0.7
32.7
"Phyaically Counted
^Estimated
-------
4.3.3. References for Section 4.3
1. Murphy, J. R., M. R. Smith, and C. H. Viens. Hydrocrack-
ing and Fluid Catalytic Cracking of Gas Oils in Today's
Refinery. In: Proceedings, API Div. of Refining, 50:
923-940, 1970.
2. Cantrell, A. Annual Refining Survey. Oil and Gas J.,
77(13): 122-156, March 26, 1979.
3. Gary, James. H., and Glenn E. Handwerk. Petroleum Refining,
Technology and Economics. Dekker, New York, 1975.
4. Hemler, C. L. Developments in Fluid Catalytic Cracking.
Nafta, Zagreb, Yugoslavia. 27(4): 207-216, 1976.
5. Stover, R. D. Control of Carbon Monoxide Emissions from
FCC Units by UltraCat Regeneration. In: Proceedings
of the Ind. Proc. Des. Poll. Control, AIChE Workshop 6:
80-85, 1975.
6. Murphy, J. R., and M. Soudek. Modern FCC Units Incorporate
Many Design Advances. Oil and Gas Journal, 75(3): 70-76,
January 17, 1977.
7. Rheaume, L., et al. New FCC Catalysts Cut Energy and In-
crease Activity. Oil and Gas J., 74(20): 103-107, 110,
May 17, 1976.
8. Murphy, J. R. Catalyst and Design Spur FCC Revival. Oil
and Gas J., 68(47), November 23, 1970.
9. Arthur D. Little, Inc. Screening Study to Determine Need
for SO and Hydrocarbon NSPS for FCC Regenerators. Final
Report. EPA-450/3-77-046. Cambridge, MA, August 1976.
10. Brunch, H. R. Refiners Focus on New FCC Technology. NPRA
Question and Answer Session. Oil and Gas J., 74(11): 87-93,
March 15, 1976.
11. Davis, John C. FCC Units Get Crack Catalysts. Chem. Eng.,
84(12): 77-79, June 6, 1977.
12. Radian Corporation. Control Techniques for Volatile Organic
Emissions from Stationary Sources. Final Report. EPA-450/
2-78-022, Austin, Texas, May 1978.
13. Barbier, J. Save Energy When Making Gasoline. Hydrocarbon
Processing, 56(9): 85-96, September 1977.
227
-------
14. Hobsorv G. D., and M. L. Peniston-Bird. Modern Petro-
leum Technology. Applied Science Publishers, ltd., Essex,
England, 1975. 293 pp.
15. Kemp, Vernon E. and Owen W. Dykema. Inventory of Combus-
tion-Related Emissions from Stationary Sources. 2nd Update.
EPA-600/7-78-100. Aerospace Corporation, Environmental and
Energy Conservation Division, El Segundo, CA, June 1978.
16. Dickerman, J. C., et al. Industrial Process Profiles for
Environmental Use: Chapter 3, Petroleum Refining Industry.
Final Report. EPA-600/2-77-023C, Radian Corporation,
Austin, Texas, January 1977.
17. Environmental Protection Agency. Compilation of Air Pol-
lutant Emission Factors. Third Edition, Supplement No. 8.
AP-42. Research Triangle Park, N.C., May 1978.
18. Environmental Protection Agency. Development of Petroleum
Refinery Plot Plans. EPA-450/3-78-025. Research Triangle
Park, North Carolina, 1978.
19. Air Pollution Control District, County of Los Angeles;
State of California, Department of Public Health; U.S.
Department of Health, Education and Welfare; Western Oil
and Gas Association. Emissions to the Atmosphere from
Eight Miscellaneous Sources in Oil Refineries. June 1958.
- 20. Unzelman, George H., and Norman H. Gerber. Hydrocracking-
Today and Tomorrow. Ethyl Corporation, Los Angeles,
California. (Undated)
21. Choudhary, N. , and D. N. Saraf. Hydrocracking: A Review.
(Ind. Eng. Chem., Prod. Res. Dev.) 14(2), 1975.
22. Chakrabarty, Shubhra, and Lalitha B. Chakrabarty. Which
Process to Maximize Middle Distillates? Chem. Eng. World
9(3): 43-45, 1974.
23. Winsor, J. Developments in Hydrocracking. Chem. Eng.
(London), 194-99, May, 1972.
24. Baral, William J., and Hal C. Huffman. Advances in Hydro-
cracking of Distillates. In: Proceedings of the 8th World
Petrol. Congr., 4: 19-27, 1971.
25. Murphy, J. R., M. R. Smith, and C. H. Viens. Hydrocracking
Vs. Cat Cracking for Gas Oils in Today's Refineries. Oil
Gas J., 68(23): 108-12, June 8, 1970.
228
-------
26. Murphy, J. R., M. R. Smith, and C. H. Vrens. Hydrocrack-
ing and Fluid Catalytic Cracking of Gas Oils in Today's
Refinery. In: Proceedings of the API Div. of Refining, 50
923-940, 1970.
27. Ward, J. W., A. D Reichle, and J. Sosnowski. Catalyst
Advances Open Doors for Hydrocracking. Oil and Gas J.,
71(22), 1973.
28. Hydrocarbon Processing Refining Process Handbook. Hydro-
carbon Processing, 57(9), September, 1973.
229
-------
4.4 r HYDROPROCESSING - c_
Hydroprocessing refers to those processes in which
hydrogen is mixed with a variety of feedstocks-and passed
over a catalyst at elevated temperature and pressure. The
hydrogen reacts with sulfur- and nitrogen-containing compounds
in the feedstock to form hydrogen sulfide and ammonia. Heavy
metals, oxygen and halides are also removed via hydroprocessing.
Hydroprocessing may also be used to stabilize unsaturated hydro-
carbons such as olefins by converting them to saturated
materials.
Hydroprocessing operations may be divided into three
categories, based on the severity of the process: 1) hydro-
cracking, in which 50 percent or more of the feed is reduced in
molecular weight; 2) hydrorefining, in which 10 percent or less
of the feed is reduced in molecular weight; and 3) hydrotreat-
ing, in which essentially no reduction in'molecular weight
occurs.
Hydrocracking is discussed in detail in Section 4.3.2.
The following sections describe hydrorefining and hydrotreating
processes.
4.4.1 Hydrorefining
Hydrorefining includes processes in which heavier dis-
tillation cuts (e.g. residuals, heavy gas oils, topped crudes,
or middle distillates) are refined or cleaned for further process-
ing, blending, or direct use. Pretreatment of light and heavy
feedstocks for catalytic cracking or reforming is also included
in this --category. A number of similar processes^are available
from various licensors which reduce the sulfur, nitrogen,
230
-------
or raeta.j~ content of heavy feedstocks. These a-reL discussed in
the following paragraphs.
4.4.1.1 Process Description and Technology--
Hydrodesulfurization of Residual Oil--The purpose of
hydrodesulfurization of residual oils is to remove contaminating
sulfur, nitrogen, and metals and thus upgrade the residual feed-
stock, for use in fuel oil production or catalytic cracking. Use
of residual hydrodesulfurization has historically been limited
due to the high rate of hydrogen consumption and the short
catalyst life. Table F4-32 lists licensors currently offering
residual hydrodesulfurization processes and the number of each
type of unit in commercial operation. Figure F4-12 is a
schematic diagram for residual hydrodesulfurization. The basic
process involves heating a mixture of feedstock and hydrogen and
passing the mixture through a reactor containing a catalyst bed.
The reactions occur in the liquid phase. Both fixed and ebul-
lient catalyst beds are used. In the reactor, sulfur in the feed
is converted to H2S; nitrogen is converted to NH3; and metal.
contaminants are deposited on the catalyst. After products
from the reactor are cooled, hydrogen and HaS are flashed in a
series of high and low pressure separators. The hydrogen is
recycled and the HaS is recovered for further processing. The
remaining oil product is steam-stripped to remove residual H2S,
and may pass through a fractionator for removal of light
hydrocarbons. The desulfurized residuum may be blended into
fuel oil or further processed.
Feed to the residual hydrodesulfurization unit is
usually residuum with boiling points ranging from 650-850°F and
a sulfur- content higher than the parent crude oil. Hydrogen is
added at- the rate of 400-700 scf/bbl. Although a variety of
231
-------
TABLE F4-32. LICENSORS OF RESIDUAL HYDRODESULFURIZATION PROCESSES3
Name of Process
Licensor(s)
No. Units
Operating
Commercially
Total
Unit
Capaclty
K>
LO
Fuel Hydrodesulfurization
llydrodesulfurization, Residual Oil
RDS/VRDS Hydrotreating
Res id HDS
Unicracking/HDS
VGO/DAO Hydrotreating
BASF AG & IFF NA
Shell Development Co. 1
Chevron Research Co. 1
Gulf Research & Development 5
Company
Union Oil Co. of California 1
Chevron Research Co. 24
NA
6,800 tpd
23,000 b/cd
NA
60,000 b/sd
845,000 b/d
tfeedstockls are primarily atmospheric or vacuum residual; however, some processes also use whol/e
crude or lighter gas oils as well.
-------
ro
CO
HYOROGEN ?
STARI
H-S SCRUBBER
** LPG OFF-GAS
LOW-PRESSURE
SEPARATORS
PRODUCT TO
FRACTIONATOR
HIGH-PRESSURE
SEPARATORS
70-1531-1
Figure F4-12. Residual hydrodesulfurization.
-------
proprietary catalysts are in use, the most commonly used are
cobalt-molybdenum or cobalt-nickel based.
The sulfur content of the desulfurized residuum varies
with the feedstock sulfur content and the severity of.the desul-
furization process. Table F4-33 shows typical product yields
and properties, investment, utility requirements and catalyst
costs as reported by Gulf Research and Development Co. for the
Resid HDS process.
Process equipment - The major pieces of process equip-
ment used in hydrodesulfurization processes include a process
heater, a fixed-bed or ebullient catalytic reactor, separator
drums, an absorber/scrubber for HzS, and a steam stripper. All
of these pieces of equipment have been discussed in previous
sections of this appendix.
Hydrodesulfurization of Heavy Gas Oils and Middle
Distillates--Heavy gas oils and middle distillates are hydrode-
sulfurized to remove sulfur, nitrogen and metallic compounds.
The process is used extensively to produce high-quality, low
sulfur kerosine and light gas oils for the production of jet
fuels, diesel fuels, and heating oils. The process is also used
to treat heavy gas oils for blending or for high quality cata-
lytic cracking feed.
Hydrodesulfurization is particularly important in
pretreating catalytic cracker feeds. Removal of sulfur and
other impurities results in reduced catalyst poisoning, longer
catalyst life, better cracking selectivity and better product
quality.. Additionally, lower-sulfur feed results in less sulfur
deposition on the catalyst; thus, catalyst regeneration produces
less SOa-emissions.
234
-------
TABLE F4-33. RESIDUAL HYDRODESULFURIZATION
Kuwait Atm Resid.
Mild
3.1
0.6
2.6
8.5
90.1
-
90
16.6
3.8
60
Moderate
3.7
0.9
3.8
-
-
98.0
640
-
Severe
3.9
1.1
4.6
-
-
97.4
720
Yields - Average for run:
Charge (650°F+)
Gravity, "API
Sulfur, wt %
Ni -t- V, ppm
HDS Severity
Product Yield
H2S, wt %
Ci-Cit, wt %
C5-375°F Naphtha, vol %
375-600°F Distillate, vol %
650°F Fuel Oil, vol %
375°F Fuel Oil, vol %
Chemical H2 cons., scf/bbl
Economics - For processing 50,000 bpsd, 650°F +• Kuwait -
Atm tower bottoms, 2 cycles per year:
HDS Severity Mild Moderate Severe
Investment
(U.S. Gulf Coast, includes
catalyst, excludes hydrogen
and sulfur) c per bpsd 704 937 1166
Utilities (per bbl feed)
Powe r, kwh 4.7 5.5 6.4
Steam (600 psig, 750°F) Ib 97 17
Steam (50 psig, sat'd), Ib 26 39 51
Fuel, heat released, 103 Btu 76 76 81
Cooling water (20°F rise), gal 200 230 360
Condensate (deaerated), gal 1.3 1.7 2.2
Catalyst Cost (c per bbl) 14 36 62
Source: Reference 8
235
-------
i-" A number of proprietary hydrodesulfurteation processes
are currently in use. Table F4-34 lists major processes, their
frequencies, capacities, and licensors.
Feed to a typical hydrodesulfurization unit may be
kerosine, light gas oil, distillate oil, or heavy gas oil. Hydro-
gen requirements vary with feed type and the severity of the
desulfurization process. Kerosine hydrodesulfurization requires
about 400 scf hydrogen per barrel feed while heavy gas oils
require up to 1700 scf/bbl.
A simplified process diagram of a typical heavy gas
oil hydrodesulfurization unit is shown in Figure F4-13. The oil
is mixed with make-up and recycle hydrogen, heated, and charged
to a fixed bed reactor containing a non-noble metal catalyst
(generally nickel or cobalt-molybdenum). The reactions occur
essentially in the liquid phase and sulfur and nitrogen in the
feed react with hydrogen to form H2S and NH3. A hydrogen-rich
stream is flashed from the reactor product in a high pressure
separator and returned to the reactor. Reactor product flows to
a low pressure separator where most of the H2S, NH3, and light
ends are recovered. The oil product is then steam-stripped or
fractionated to remove the remaining impurities.
Product yield varies with the type of feed processed.
Licensors of most processes report sulfur-removing capabilities
of 90 percent or greater, with product yields ranging from 98-100
volume precent of charge.2
236
-------
TABLE F4-34. HEAVY GAS OIL AND MIDDLE DISTILLATE
HYDRODESULFURIZATION PROCESSES
Name of Process
GO-f ining
Gulfing
Licensor
Exxon Research &
Engineering Company
Gulf Research &
No. Units
Operating
Commercially
23
6
Total
Capacity
106 b/sd
NA
Hydrodesulfurization
Trickle Flow
Hydrofining
Hydrofining
RCD Unibon
Development Co. &
Houdry Div. of Air
Products & Chemicals,
Inc.
Shell Development Co.
BP Trading Ltd.
Exxon Research &
Engineering Co.
UOP Process Div.
of UOP Inc.
91
49
225
NA
185,000 t/d
720,000 b/sd
A x 106 b/sd
NA
NA - Not available
Source: Reference 2
-------
N>
U)
oo
HYDROGEN
MAKE-UP
RECYCLE
c — "•
1
A
— ->
~l J
RFACTOR
Hintt-PRESSUIlE
SEPARATOR
FUEL GAS
SCRUBBER
STEAM STRIPPER
DESUI.FURIZED
PRODUCT
70-1W7-1
Figure FA-13. Heavy gas oil hydrodesulfurination,
-------
^Process conditions, fuel and utility requirements -
Temperature: 390-800°F
Pressure: 500-800 psi
Electricity: 19-365 kWh/bbl
Heater Fuel: 0-70,000 Btu/bbl
Steam: 1-10 Ib/bbl
Cooling Water: 160 gal/bbl
Process equipment - The major pieces of process equip-
ment used in heavy gas oil hydrodesulfurization are
process heaters, a fixed bed catalytic reactor, high-
and low-pressure separator drums, and a steam stripping
column. All have been previously described in this
appendix.
4.4.1.2 Atmospheric Emissions--
Emissions from hydrorefining processes include:
• Emissions during periodic catalyst regeneration,
• Process heater flue gas, and
• Fugitive emissions.
Emissions from Catalyst Regeneration Operations--
During the hydrorefining process, deposits form on the catalyst
which must be removed to restore catalyst activity and selec-
tivity. This regeneration procedure involves burning the
deposits by admitting air into the reactors.
During regeneration, large quantities of carbon
monoxide-and other pollutants may be released. However,
239
-------
regeneration may be required only after several "months or even
years of operation. Hence, total averaged emissions from this
source are generally insignificant.
Process Heater Flue Gas--Emission factors for various
pollutants from fired heaters are given in Table F4-35. A
typical hydrorefiner will require a process heater to heat the
combined hydrocarbon-hydrogen feed to the reactor. Heat input
requirements vary, and may be as high as 100,000 Btu per barrel
of oil feed.
Total estimated emissions from typical gas oil hydro-
desulfurization heaters are also given in Tabel F4-35. These
figures are given as pounds of pollutant per thousand barrels of
feed using a representative heat input of 60,000 Btu per barrel
feed.
Fugitive Emissions—Fugitive emissions can occur from
a variety of sources including valves, pumps, compressors,
flanges and other fittings, relief valves, and drains. Hydro-
carbon emission rates for each individual source are quite low.
However, total fugitive emissions are significant because of the
large number of such sources within the refinery.
Total estimated fugitive non-methane hydrocarbon emis-
sions from a typical gas oil hydrodesulfurization unit are given
in Table F4-36. The listed emission factors were determined as
a part of this program. Additional information on these emis-
sion factors, including a complete discussion on their deriva-
tion and the confidence intervals for each source "category, are
contained in Appendices B and C of this report.
".Estimates for the number of sources within each source
category were developed from:
240
-------
F4-35. TYPICAL EMISSIONS FROM GAS-OIL HYDRO-
DESULFURIZATION UNIT PROCESS HEATERS
a
EPA Emission Factor
(lb/103 gal-oil fired)
(lb/106 scf-gas fired)
Total Emissions
(lb/103 bbl of
"fresh feed)
Oil Fired Heaters
Particulates
- Distillate oil
- Residual oil
Grade 4
Grade 5
Grade 6
Sulfur Dioxide
- Distillate oil
- Residual oil
Sulfur Trioxide0
Carbon Monoxide
Hydrocarbons (as
Nitrogen Oxides
(as N02)
- Distillate oil
- Residual oile
7
10
10(S)+3
142(S)
157(S)
2(S)
5
1
22
22+400(N)"
0.86
3.0
4.3
4.3(S)+1.3
60.9(S)
67.3(S)
0.86(S)
- 2.1
0.43
9.4
9.4+17100"
Gas Fired Heaters
Particulates
Sulfur Oxides (as S02)
Carbon Monoxide
Hydrocarbons (as CELJ
Nitrogen Oxides (as N02)
5-15
0.6
17
3
120-230
0.29-0.86
0.034
0.97
0.17
6.86-13.1
*Source: Reference 6
Based on a heat input of 60,000 Btu/bbl of fresh feed with the following
fuel heating values: Oil - 140,000 Btu/gal; Gas - 1050 Btu/scf.
CS • vt Z sulfur in the oil
Improper combustion may cause a significant increase in emissions
^se this emission factor for residual oils with less than 0,32 QJ<.5) nitro-
gen content. For oil with higher nitrogen content QJ>0,51, use emission
factor of 120 lb/103 gal
Basea on sulfur content of 2000 gr/106 scf
241
-------
TABLE F4-36.
ESTIMATED FUGITIVE NON-METHANE HYDROCARBON EMISSIONS FROM
A TYPICAL GAS OIL HYDRODESULFURIZATION UNIT
N)
£-
K>
\
Emissions
Source Type
Valves
Open-End
(Sample)
Valves
Pumps (Pump
Seals)
Drains
Flanges &
Fittings
Process
Stream Service
Classification
Gas/Vapor
Light Liquid
(VP > 0.1 psla 9 100'P)
Heavy Liquid
(VP < 0.1 psla 9 100'P)
Hydrogen Service
Total
All
Light Liquid
(VP > 0.1 p«l« 9 100'F)
Heavy Liquid
(VP < 0.1 pala 9 100'P)
Total
All
All
Relief Valves All
Compressors
(Compressor
'Seals)
Hydrocarbon
Hydrogen
Total
Number of Sources in
Counts or Estimates
From Radian Study
235
208
102
101
645a
, -
6( 9)
4( 5)
10(14)fl
24a
2743a
6C
0(0)
3(6)
3(6)a
Process Unit
Counts or Estimates
From PES Study
205
244
97
164
710C
L
16b
5( 7)
2( 3)
7(10)b
-
2350C
-
0(0)
3(6)
3(,6)b
Source
Emission
Factor, Ib/hr
0.059
0.024
0.0005
0.018
0.005
0.25
0.046
0.070
0.00056
0.19
1.4
0.11
Estimated Total
Emissions,
Ib/hr
13.9 - 12.1
4.99 - 5.86
0.049- 0.051
1.82 - 2.95
20.8 - 21.0
0.080
1.75 - 2.25
0.14 - 0.23
1.89 - 2.48
1.68
1.44 - 1.54
1.14
0.0
0.66
, .' 0.66
/•''
27.7 - 28.6
Physically Counted
Counted From Flow Diagrams
CEstlraated
Reference 7
-------
~~« Actual field source counts conduc&ed during the
course of this program, and
• Counts contained in The Development of Petroleum
Refinery Plot Plans, EPA publication number
450/3-78-025.7
The methods used to develop estimates for the number
of each source type have been previously described in Section
4.1.1.3 of this appendix.
The source counts listed in Table F4-36 refer to those
sources located within the battery limits of the process. That
is, equipment located in tankage or transfer lines is not
included in these source counts.
The number of sources per unit depends on a variety of
factors including processing complexity, process type, and
processing flexibility. The source counts given in Table F4-36
are not necessarily representative of all or even the majority
of hydrocracking units.
The estimated composition of fugitive non-methane
hydrocarbon emissions from a gas oil hydrodesulfurization unit
is given in Table F4-37. These estimates indicate the types of
hydrocarbons contributed by sources on a number of streams within
the process unit. Additional information about these estimates
may be found in Section 2.4 of Appendix D.
4.4.2 Hydro treating
Hydro treating processes are generally less severe than
hydrorefining processes. Feedstocks for these operations include
243
-------
TABLE F4-37.
ESTIMATED COMPOSITION OF FUGITIVE NON-METHANE EMISSIONS
FROM A GAS OIL HYDRODESULFURIZATION UNIT
NJ
Stream
Estimated percentage of emissions
attributed to each stream - Z
Weighted contribution of each
component to unit emissions — ppmw
Benzene
Toluene
Ethylbenzene
Xylenes
Other Alkylbenzenes
Naphthalene
Anthracene
Biphenyl
Other Polynuclear aroroatics
n-Hexane
Other Alkanes
Olefins
Cyclo Alkanes
Hydrogen
Gaa Oil
22
0
1
1
5
8
6
2
2
146
0
208756
0
11000
0
Desulfurlzed
Gaa Oil
22
0
1
1
5
81
6
2
2
146
0
208756
0
' 11000
0
H2 Recycle
Gas
56
0
0
0
0
0
0
0
0
0
0
364000
0
0
196000
Totals
100Z
0
2
2
10
162
12
4
4
292
0
781512
0
, .' 22000 '
196000 '
1000000
-------
naphtha, ..straight run distillate, olefins, aromat-ics and
processed lube oil. As in hydrorefining, the hydrotreating pro-
cess serves to remove sulfur, nitrogen, and metallic compounds
from the feed. Hydrotreating is also used to saturate, and
thus stabilize, olefins and aromatics, and to polish and dewax
lube oil stocks.
The following sections describe three types of hydro-
treating processes in current commercial use: naphtha/SR
distillate desulfurization; aromatic saturation, and lube oil
hydrotreating.
4.4.2.1 Process Description and Technology--
Light Hydrocarbon Hydrodesulfurization--The light
hydrocarbon hydrodesulfurization unit is used to desulfurize
and denitrogenate the naphtha and straight run (SR) distillate
streams from the crude distillation unit. Both sulfur and
ammonia must be removed because the light hydrocarbon streams
are primary feeds for the isommerization and catalytic reform-
ing units.
A number of proprietary processes in commercial use
are listed in Table F4-38. Information describing unit fre-
quency, capacity, and licensor is also included.
Feed to the light distillate hydrodesulfurization unit
is typically a mixed stream of light hydrocarbons (sour naphtha,
kerosine, diesel oil, and other light straight-run or cracked
distillates) with a boiling range of 100-430°F. Make-up hydrogen
is usually necessary to produce the desired degree of hydrogena-
tion, although one licensor offers a process based on simultan-
eous dehydrogenation and desulfurization in which"no make-up
hydrogen is required.2
245
-------
TABLE F4-38. LIGHT DISTILLATE IIYDRODESULFURIZATION PROCESS
(Ti
Name of Process
Autof ining
Distillate Hydrodcsulfurization
Hydrodesulfurizatlon,
Vapor Phase
Hydrof ining
Hydrof ining
Unionf ining
Licensor
BP Trading Ltd.
Institut Francais
du Petrole
Shell Development Co.
BP Trading Ltd.
Exxon Research &
Engineering Co.
Union Oil of
California
Number of
Commercial Units Capacity
4 9,800 b/sd
50 NA
72 185,000 tpd
49 720,000 b/sd
225 4 x 10° b/sd
75 NA
NA = Not available
Source: Reference 2
-------
~~A schematic diagram of light distilla-te^ hydrodesul-
furization is shown in Figure F4-14. Gaseous-phase feedstock
is mixed with a hydrogen-rich gas and heated to reaction temper-
ature. The mixture is then passed through a fixed-bed reactor
containing a non-noble metal catalyst, usually cobalt-molybdenum
or cobalt-nickel. In the reactor, organic sulfur and nitrogen
compounds react with hydrogen to form HaS and NHs. Some crack-
ing of naphthas into lighter fractions may occur as a side
reaction.
The hot effluent is cooled in heat exchangers and is
then sent to a high pressure separator where hydrogen is
flashed and recycled to the feed stream. The liquid from the
separator is sent to a fractionator where FUS, NH3, and light
hydrocarbons boil off and are sent to an amine unit for removal
of the acid gases. The remaining liquid product is further
fractionated or continues to the reformer or isomerization
units. Product yields of 99 percent or greater are reported
for most commercial processes.
Process conditions, fuel, and utility requirements -
The operating conditions and fuel/utility requirements
of light hydrocarbon hydrodesulfurization units vary
depending on feedstock composition. However, the
following operating ranges are typical for most
commercial processes:
Temperature: 600-800°F
Pressure: 300-1000 psi
Electricity: 2.6 kWh/bbl
Heater Fuel: 36,000-75,000 Btu/bbl
Cooling Water.- 264 gal/bbl
- Steam: 30-90 Ib/bbl if steam stripper^used, 5 Ib/bbl
without steam stripper
247
-------
NJ
*~
00
HYDROGEN.
MAKE-UP '
FEEO 1
H.P. PURGE
OFF-GAS
HYDROGEN RECYCLE
HIGH-PRESSURE
SEPARATOR
REACTOR
FRACT10NATOR
LIGHT
HYDROCARBONS
DESULFURIZED
PRODUCT
70-1530-1
Figure F4-14. Light distillate hydrodesulfurization.
Source: Reference 9
-------
'•• ~ Process equipment - The major pieces'Tbf^process
equipment in light hydrocarbon hydrodesulfurization
units include process heaters, a fixed bed catalytic
reactor, heat exchangers, a high pressure separator
drum, a fractionation column, and (optionally) a
steam stripper. This equipment is described in
detail in previous sections of this appendix.
Olefin/Aromatics Saturation—The purpose of olefin/
aromatic saturation is to stabilize these compounds by converting
aromatics to paraffins and by de-gumming olefins and diolefins.
The stabilized products may be used as feedstocks and blending
stocks for gasoline, aviation fuels, and turbine fuels. Some
processes produce high-purity aromatic streams which may be sent
to the aromatics extraction unit.
Several commercial processes are available for
saturating olefins and aromatics. These processes are listed
in Table F4-39.
Feed to the saturation unit may be pyrolysis naphtha,
turbine fuel stocks, olefins, diolefins, and other petroleum
streams with high aromatic content. Hydrogen requirements vary
with the type of feedstock charged, but generally range between
450-700 scf/bbl charge.
Figure F4-L5 is a simplified process diagram for
olefin/aromatic saturation. Feed to the unit is combined with
make-up and recycled hydrogen and heated in a fired heater. The
heated mixture then enters a fixed-bed reactor containing a
platinum-, palladium- or nickel-alumina catalyst. Reactor
effluent -is cooled and separated into liquid and vapor phases
in the separator. The resulting vapor phase is recycled to the
249
-------
TABLE F4-39. PROCESSES FOR OLEFIN/AROMATICS SATURATION
Cn
O
Number of Total
Name of Process Licensor Commercial Units Capacity
Arosat C. E. Lummus 1
DPG Hydrogenation C. E. I.ummus 23
HPN Engelhard Industries 25
Unisar Union Oil of California 6
NA
160,000 b/sd
NA
50,000 b/d
NA = Not available
Source: Reference 2
-------
Ln
HYDROGEN
MAKE-UP
COMPRESSOR
LIGHT
ENDS
RECYCLE HYDROGEN
FEED
REACTOR
SATURATED
PRODUCT
70-1529-1
Figure F4-15. Saturation of aromatics/olefins.
Source.- Reference 2, p. 155
-------
reactor - a*i~d the saturated liquid product is sent-to the stripper
to remove light ends. After stripping, the saturated product
may be sent to further processing or blending operations.
Product yields range from 100-104 volume percent feed
depending on the feedstock charged. The relative amounts of
paraffins, naphthas, and aromatics in the charge also depend on
the type of charge.
Process conditions, fuel, and utility requirements -
Temperature: 480-660°F
Pressure: 100-1500 psi
Electricity: 0.5-2.5 kWh/bbl feed
Steam: 12-35 Ib/bbl
Fuel: 103 Btu/bbl
Cooling Water: 120-680 gal/bbl
Process equipment - The major equipment items in an
olefin/aromatic saturation unit are a process heater,
a fixed-bed catalytic reactor, a pressurized separator
drum, and a steam stripping column. All have been
described in detail in previous sections of this
appendix.
Lube and Wax Hydrotreating--Lube oil and wax stocks
are hydrotreated to improve product quality in several ways:
viscosity index improvement, desulfurization, denitrogenation,
demetallization, removal of gum-forming compounds, and color
improvement. The products of this unit are primarily finished
single- or multi-grade lube oils, but gasoline, naphtha, kero-
sine, furnace oil, and waxes are valuable by-products.
-Table F4-40 lists major commercial processes for lube,
fuel, and" specialty oil hydrotreating and dewaxing.
252
-------
TABLE F4-40. HYDROTREATING PROCESSES FOR OILS AND WAXES
Ui
OJ
Name of Process
Distillate Dewaxing
Ferrofining
Hydrogen Finishing
Lube, Wax, and Specialty
Oil Hydrotreating
Lube Oil HydroLreating
Wax Hydrofinishing
Licensor
Number of
Commercial Units
Mobil Oil Corporation
BP Trading Ltd.
Texaco Development Corp.
Institut Francais due Petrole,
and Total Companie Francais de
Raffinage
Gulf Research and Development Co.
BP Trading Ltd.
1
NA
NA
10
4
I
Total
Capacity
NA
54,200 b/sd
NA
NA
NA
50 tons/day
NA = Not available
Source: Reference 2
-------
• ~ Feed to a lube/wax hydrotreating unit^may be either
solvent-refined lube oils and waxes, or raw distillates and
deasphalted oils. Hydrogen requirements vary from 100-200 scf/
bbl charge.
A simplified process diagram for lube oil/wax hydro-
treating is shown in Figure F4-16. The oil feed is mixed with
make-up hydrogen and charged to a fixed bed reactor. The
catalyst is usually either cobalt- or nickel-molybdenum on an
alumina base. Reactor effluent flows through high and low
pressure separators to remove hydrogen for recycle and light
ends, respectively. The product is then charged to a double
stage column where it is steam-stripped and vacuum dried.
Product yields are very high, varying with the
severity of the process. Table F4-41 lists product yields from
a typical lube oil/wax hydrotreating unit.
TABLE F4-41. PRODUCT YIELDS FROM LUBE OIL/WAS HYDROTREATING
Product
(Volume % Charge)
€3-65 Hydrocarbons
Cs Furnace Oil
Light Lube
Light Neutral
Heavy Neutral
Bright Stock
Slack Wax (4% Oil Content)
TOTAL
Severity
Medium
1.5
18.9
4.6
24.0
33.9
10.6
13.0
106.5
High
3.0
57.0
14.4
20.5
11.6
-
8.0
114.5
Source: Reference 2, p. 143.
254
-------
to
ui
L/l
HYDROGEN
MAKE-UP '
RECYCLE HYDROGEN
LIGHT
ENDS
COMPRESSOR
FEED ?—eJ fr
DISTILLATE
REACTOR
1
HIGH-PRESSURE
SEPARATOR
LOW-PRESSURE
SEPARATOR
STRIPPER/VACUUM
DRYER
STtAH
PRODUCT
70-1528-1
Figure F4-16. Lube oil hydrotreating.
Source: Reference 2, p. 142
-------
^'Process conditions, fuel, and utility-requirements -
Reactor Temperature: 600-750°F
Reactor Pressure: 500-700 psi
Electricity: 2.5 kWh/bbl charge
Steam: 15-30 Ib/bbl charge
Heater Fuel: 35,000-140,000 Btu/bbl charge
Process equipment - The major pieces of process equip-
ment in a lube oil/wax hydrotreating unit are a
process heater, fixed-bed reactor, high and low
pressure separators, and a double stage steam
stripping/vacuum drying column. All of this equipment
has been described in previous sections of this
appendix.
A.4.2.2 Atmospheric Emissions--
Emissions from hydrotreating processes include:
• Emissions during periodic catalyst regeneration,
• Process heater flue gas, and
• Fugitive emissions.
Emissions from Catalyst Regeneration Operations--
During hydrotreating processes, deposits form on the catalyst
which must be removed to restore catalyst activity and selec-
tivity. The regeneration procedure involves burning the deposits
by admitting air into the reactors.
During regeneration, large quantities of carbon
monoxide-and other pollutants may be released. However,
256
-------
regeneration may be required only after several"months or even
years of operation. Hence, total averaged emissions from this
source are generally insignificant.
Process Heater Flue Gas — Emission factors for various
pollutants from fired heaters are given in Table F4-42. A
typical hydrotreater will require a process heater to heat the
combined hydrocarbon-hydrogen feed to the reactor. Heat input
requirements vary, ranging from 35,000-140,000 Btu per barrel
of oil feed.
Total estimated emissions from typical hydrotreater
heaters are also given in Table F4-42. These figures are given
as pounds of pollutant per thousand barrels of feed using a
total required heat input of 75,000 Btu per barrel feed.
Fugitive Emissions—Fugitive emissions can occur from
a variety of sources including valves, pumps, compressors,
flanges and other fittings, relief valves, and drains. Hydro-
carbon emission rates for each individual source are quite low.
However, total fugitive emissions are significant because of
the large number of such sources within the refinery.
Total estimated fugitive non-methane hydrocarbon emis-
sions from a typical hydrotreating unit are given in Table F4-43.
The listed emission factors were determined as a part of this
program. Additional information on these emission factors, in-
cluding a complete discussion on their derivation and the con-
fidence intervals for each source category, are contained in
Appendices B and C of this report.
Estimates for the number of sources within each source
categoryjwere developed from:
257
-------
^ABLE F4-42.
TYPICAL EMISSIONS FROM HYDEQTREATING
UNIT PROCESS HEATERS
EPA Emission Factor
(lb/103 gal-oil fired)
(Ib/lO* scf-gas fired)
Total Emissions
(lb/103 bbl of
fresh feed)
Oil Fired Heaters
Participates
- Distillate oil
- Residual oil
Grade 4
Grade 5
Grade 6
Sulfur Dioxide0
- Distillate oil
- Residual oil
Sulfur Trioxidec
Carbon Monoxide
Hydrocarbons (as
Nitrogen Oxides
(as N02)
- Distillate oil
- Residual oile
7
10
10(S)+3
142 (S)
157(S)
2(S)
5
1
22
22+400 (N)'
1.1
3.8
5.4
5.4(S)+1.6
76.1(S)
84.1(3)
l.l(S)
2.7
0.54
12
12+214(N)^
Gas Fired Heaters
Participates
Sulfur Oxides (as S02)
Carbon Monoxide
Hydrocarbons (as CEO
Nitrogen Oxides (as N02)
5-15
0.6
17
3
120-230
0.36-1.1
0.043
1.2
0.21
8.6-16.4
Source: Reference 6
Based on a heat input of 75,000 Btu/bbl of fresh feed with the following
fuel heating values: Oil - 140,000 Btu/gal; Gas - 1050 Btu/scf.
CS • wt Z sulfur in the oil
Improper combustion may cause a significant increase In emissions
^se this emission factor for residual oils with less than 0.52 0«.5) nitro-
gen content. For oil with higher nitrogen content QT>0,5}, use emission
factor of 120 lb/103 gal
Based on sulfur content of 2000 gr/106 scf
258
-------
TABLE F4-43.
ESTIMATED FUGITIVE NON-METHANE HYDROCARBON EMISSIONS FROM
A TYPICAL HYDROTREATING UNIT
I , f Process
Emissions Stream Service
Source Type Classification
Valves Gas/Vapor
Light Liquid
(V? > 0.1 P«t« 9 100'F)
Heavy Liquid
(VP < 0.1 ps la 8 100V)
Hydrogen Service
Total
Open-End
(Sample) All
Valves
Putt pa (Pump Light Liquid
Seals) (vp > o.l p>u 9 100T)
Heavy Liquid
(vp f 0.1 p»i« t 100*P)
Totol
Drains All
Flanges &
Fittings All
Relief Valvea All
Compressors Hydrocarbon
(Compressor Hydrogen
Seals) Total
"Physically Counted
Counted Fron Flow Diagrams
C Estimated
Reference 7
Number of Sources In Process Unit
Counts or Estimates Counts or Estimates
From Radian Study From PKS Study
235 226 - 389
208 378 - 648
102 0
101 181 - 312
645a 785 - 1349C
17 - 29b
6( 9) 8(11)-16(22)
4( 5) 0
10(14)a 8(ll)-16(22)t>
2/,a
2743a 2585 - 4465<=
6C
0(0) 0(0)
3(6) 3(6)
3T6)° 3(6)b
Source
Emission
Factor, Ib/hr
0.009
0.024
0.0005
0.018
0.005
0.25
0.046
0.070
0.00056
0.19
1.4
0.11
'
Estimated Total
Emissions,
Ib/hr
13.3 - 23.0 '
4.99 - 15.6
0.0 - 0.051
1.82 - 5.62
20.1 - 44.3
0.085- 0.145
2.25 - 5.50
0.0 - 0,23
2.25 - 5.73
1.68
1.44 - 2.50
1.14
0.0
0.66
0.66
27.4 - 56I01
-------
-'• Actual field source counts conduc&ed during the
course of this program, and
• Counts contained in The Development of Petroleum
Refinery Plot Plans, EPA publication number
450/3-78-025.7
The methods used to develop estimates for the number
of each source type have been previously described in Section
4.1.1.3 of this appendix.
The source counts listed in Table F4-43 refer to those
sources located within the battery limits of the process. That
is, equipment located in tankage or transfer lines is not
included in these source counts.
The number of sources per unit depends on a variety
of factors including processing complexity, process type, and
processing flexibility. The source counts given in Table F4-43
are not necessarily representative of all or even the majority
of hydrotreating units.
The estimated composition of fugitive non-methane emis-
sions from a hydrotreating unit is given in Table F4-44. These
estimates indicate the types of hydrocarbons contributed by
sources on a number of process streams found within the unit.
Additional information about these estimates may"be found in
Section 2.4 of Appendix D.
260
-------
TABLE',F4-44.
N5
cr>
ESTIMATED COMPOSITION OF FUGITIVE NON-METHANE
EMISSIONS FROM A HYDROTREATING UNIT
Stream
Estimated percentage of emissions
attributed to each stream - X
Weighted contribution of each
component to unit emissions - ppmw
Benzene
Toluene
Ethylbenzene
Xylenes
Other Alkylbenzenes
Naphthalene
Anthracene
Biphenyl
Other Polynuclear aroma tics
n-Hexane
; Other Alkanea ,
Olefina
Cyclo Alkanea
Hydrogen
Straight Run
Naphtha
47
119
1232
417
763
7792
688
2
295
7042
18254
234817
0
198579
0
Desulfurized
Naphtha
47
119
1232
417
763
7792
688
2
295
7042
18254
, 234817
0
198579
0
Hj Recycle
Gas Totals
6
0
0
0
0
0
0
0
0
0
0
39000
0
0
21000
100X
238
2464
834
1526
15584
1376
4
590
14084
36508
i .' 508634
0
397158
21000
1000000
-------
4.4.3 References for Section 4.4
Dickennan, J. C., et al. Industrial Process Profiles for
Environmental Use. U. S. Environmental Protection Agency.
EPA/2-77-023C. Research Triangle Park, North Carolina,
January 1977.
Hydrocarbon Processing Refining Process Handbook. Hydro-
carbon Proc., 57(9), September 1978.
Annual Refining Survey. Oil and Gas J., 77(13): 122-156,
March 26, 1979.
McCulloch, D. C. Feed Hydrotreating Improves FCCU Per-
formance. Oil and Gas J., 73(20): 53-58, June 21, 1975.
Hobson, G. D., and W. Pohl. Modern Petroleum Technology,
4th Ed. Applied Science Publishers Ltd., Essex, England,
1975.
Environmental Protection Agency. Compilation of Air Pol-
lutant Emission Factors. Third Edition, Supplement No. 8.
AP-42. Research Triangle Park, N.C., May 1978.
Environmental Protection Agency. Development of Petroleum
Refinery Plot Plans. EPA-450/3-78-025. Research Triangle
Park, North Carolina, 1978.
Ondish, G. F., et al. Save Crude: Feed Resid to FCC.
Hydrocarbon Processing, 55(7): 105-108. July 1976.
Refinery Process Handbook. Hydrocarbon Processing, 57(9),
September 1978.
-------
4.5 . CONVERSION PROCESSES
Conversion processes may be characterized as those
processes which utilize catalyzed chemical reactions to upgrade
certain refinery streams or produce valuable products from less
valuable materials. Conversion processes allow refiners to:
• Increase the octane value of gasolines
produced within certain refining operations,
• Provide flexibility to control product
distribution to meet market requirements,
• "Rebuild" certain light hydrocarbon gases
into gasoline, and
• Produce petrochemical feedstocks.
The conversion processes considered here include:
• Catalytic reforming,
• Alkylation,
• Isomerization, and
• Hydrodealkylation.
Catalytic reforming is the process of converting C6
and heavier paraffins and naphthenes into aromatics. Common
feedstocks include straight run gasoline, hydrocracker gasoline,
and coker gasoline. The feedstocks are generally hydrotreated
263
-------
to remove sulfur, nitrogen and metallic contaminants which
adversely affect the catalyst.
The products include a highly aromatic high octane
reformate, and hydrogen. The reformate can be_used as a
gasoline component, or as a source of aromatics for the produc-
tion of petrochemicals. The hydrogen can be used for hydro-
cracking or for other hydrogen consuming processes. The catalytic
reforming process is discussed in detail in Section 4.5.1.
Alkylation refers to the chemical combination of
isobutane and olefins such as propylene and butylene to produce
a high octane material in the gasoline boiling range. This
process was originally developed to provide high octane aviation
fuel and has continued as a major refining operation due to the
continued upward trend of gasoline octane requirements.
The olefin feedstocks for alkylation are generally
obtained from the catalytic cracking process. Isobutane sources
include:
• Naturally occurring isobutane in crude oil,
• Isobutane produced by hydrocracking
processes,
• Isobutane produced by the isomerization of
normal butane, and
• Butanes purchased from sources outside the
refinery.
264
-------
-• Two distinctly different types of al-kylation processes
are available; sulfuric acid catalyzed processes, and hydro-
fluoric acid catalyzed processes. Both of these processes are
discussed in detail in Section 4.5.2.
Isomerization processes are used by refiners to
convert normal paraffins into isoparaffins. The feedstocks most
commonly isomerized are butanes, pentanes, and hexanes.
Butane isomerization units are typically used in
conjunction with alkylation units as a source of isobutane. In
some cases, these units have been constructed with the alkyla-
tion process using common distillation equipment.
Pentane and pentane-hexane isomerization units serve
to produce isopentane and/or isohexane. The octane ratings of
these isoparaffins are considerably higher than their corres-
ponding normal paraffin isomers. The isoparaffins produced may
be used as high octane-low boiling range gasoline components.
This is particularly important in view of regulations limiting
the volatility, aromatics content, and lead content of gasoline.
More detailed information on isomerization is included in Sec-
tion 4.5.3.
Hydrodealkylation is used by some refiners for the
preparation of petrochemical feedstocks. The most common use
of this-.-process is the dealkylation of toluene -and the xylenes
to produce more valuable benzene. Other applications include
the production of naphthalene from alkyl-naphthalenes. Further
information on this process is contained in Section 4.5.4.
An additional conversion process utilized by some
refiners is polymerization. Polymerization processes utilize
265
-------
sulfuric- or phosphoric acid catalysts to polymerize light
olefins"such as propylene, butylene, or isobutyl/ene into gaso-
line boiling range materials. This process, however, appears
to have been replaced by alkylation as a means of producing
gasoline from light olefins. Hence, further discussion of this
process has been omitted.
4.5.1 Catalytic Reforming
Catalytic reforming is one of the most important of
all refining process operations. In catalytic reforming, rela-
tively low octane naphthas are converted to highly aromatic,
high octane gasoline blending stocks. Modern units are capable
of continuous production of reformates with octane ratings over
100 RON clear. This is particularly important in light of con-
tinuous increases in the octane ratings required of modern
gasoline.
Since its commercial introduction in 1939, reforming
capacity has continued to expand rapidly. U.S. reforming capac-
ity now (1979) exceeds 3,700,000 barrels per.day.3
4.5.1.1 Process Description--
Basically, the reforming operation consists of con-
tacting oil and hydrogen with catalyst in a series of 3 to 6
reactors. Because the overall reaction is endothermic, the
mixture must be heated in a fired heater prior to its introduc-
tion into each of the reactors.
A flow diagram for a typical semi-regenerative re-
former is given in Figure F4-17. As indicated "in the diagram,
the mixture from the final reactor is cooled and sent to a
266
-------
N3
DESIII.FURIZEO
NAPHTHA FEED
*• NF.T HYDROGEN TO RFFINERY
HYDROGEN 10
NAPHTHA IIYOROTREATER
TO
»-LIGHT-ENDS
RECOVERY
Cr RLFORHATE
70-1521-1
Figure F4-17. Process flow diagram - catalytic reforming.
Source: Reference 7, p. 165
-------
separator where hydrogen produced during the reaction is re-
moved. ^A portion of this hydrogen is recycled" Ea~ck to the re-
actors while the remainder is available for use in a. variety of
hydrogen consuming processes.
The separator liquid is sent to a stabilizer for
separation of light hydrocarbons from the reformate product.
The reformate can then be blended into gasoline or processed in
an aromatics extraction unit for chemical production.
4.5.1.2 Process Technology--
A variety of reactions occur simultaneously during the
reforming process. Some of the more important reactions are
discussed below.1'2
A. Dehydrogenation reactions
1. Dehydrogenation of alkylcyclohexanes to
aromatics.
H2 H
I I
C CH3 C
/ \ / S \
H2-C C-H H-C C-CH3
I I >. I II .-,„
H2-C C-H2 *^ H-C C-H + 3H2
Y Y
H2 H -
2. Dehydrocyclization of paraffins to
aromatics.
268
-------
H
I •-
C "~
// \
H-C C-CH3
n_C7Hl5 - * H-C S-H
C
I
H
B. Isomerization reactions
1. Isomerization of alkylcyclopentanes
followed by dehydrogenation to aromatics
H2 H2
I I
C CH3 C
/ \ / /" \
H2-C C-H H2-C C-H2
\ / _ * I I
,C— C^ H2-C C-H2
H
H-C C-H
\ /
C
I
H
I
H2
2. Isomerization of normal paraffins to
isoparaffins.
CH3
I
CH3 — CH2—CH2—CH2 —CHa >• CH3 — CH2 — CH —CH3
269
-------
C. Hydrocracking reactions
Cio Hzz + H2 *• CeHiu + CuHio
The dehydrogenation reactions are highly endothermic
(800-1000 Btu/lb) and cause a decrease in temperature as the
reaction proceeds. Hence, the reaction mixture is reheated
prior to entering each of the reactors.1
Dehydrogenation reactions are highly desirable in that
they produce aromatics. The yield of aromatics is favored by:2
• High temperature (increases the reaction rate
but adversely affects chemical equilibrium)
• Low pressure (shifts equilibrium in favor of
aromatic products)
• Low space velocity (promotes approach to
equilibrium)
• Low hydrogen to hydrocarbon mole ratios (shifts
equilibrium in favor of aromatics).
The isomerization of normal paraffins to isoparaffins
is not as important in terms of raising the product octane as
are the_reactions producing aromatics. Although the octane
numbers of isoparaffins are considerably higher than those of
normal paraffins, the equilibrium concentrations of the iso-
paraffins fall off rapidly with increasing temperature and are
quite low at the temperatures used in reforming operations.2
270
-------
^- The effect of hydrocracking reactionsJLs twofold.
Hydrocracking reactions generally improve the octane number of
the reformates by breaking up low octane paraffins. . However,
the cracked products may be light hydrocarbons, thereby reducing
the overall yield of gasoline boiling range components. Hydro-
cracking reactions are favored by high temperature, high pres-
sure, and low space velocity.
In addition to the reactions discussed above, reac-
tions which lead to the formation of coke on the catalyst are
also important. Within the reactor, small amounts of aromatic
polymers containing naphthalene, anthracene, and heavier com-
pounds are formed. Reactions involving these compounds prob-
ably lead to the formation of coke on the catalyst surface.
Since coke buildup causes a gradual decrease in catalytic ac-
tivity, the catalyst must be regenerated periodically.
The formation of coke is enhanced by high temperature
and low hydrogen partial pressure. Therefore, the choice of
reactor operating conditions is a compromise between high con-
version to aromatics with high coke deposition on the catalyst
(low pressure operation), and long catalyst on-stream times be-
tween regeneration with lower conversion to aromatics (high
pressure operation).1
Process Types — Catalytic reforming processes can be
classified by the method or frequency of cataly-st regeneration.
Currently, three regeneration schemes are available:6
• Cyclic regeneration,
• Semi-regeneration, and
271
-------
« Continuous regeneration.
Flow diagrams for reformers featuring each, of these
regeneration techniques are given in Figure F4-18.
Cyclic regeneration processes utilize an extra reac-
tor, called a swing reactor. This allows the reactors to be
regenerated, one at a time, without affecting the operation of
the unit. Units using this regeneration technique are capable
of processing poor quality feedstocks under severe conditions
because the catalyst can be regenerated frequently. Regenera-
tion frequencies of one to two reactors per day are typical.2
In semi-regenerative processes, all of the reactors
are regenerated simultaneously at the end of a cycle which may
range from 3-24 months in length.2 Since long run times are
desired, conditions which result in rapid coke formation are
avoided. Generally, this requires higher operating pressures
with a drop in gasoline yield and octane over that of cyclic
regeneration systems. However, this is compensated to a degree
because the swing reactor and the more complex piping system
needed for cyclic operations are not required.
In comparing the relative merits of these two regen-
eration schemes, the following generalizations can be made.6
These factors favor the semi-regenerative system:
• Octane numbers above 98 RON are not required.
• Minimum plant cost is desired.
• Gasoline and hydrogen have relatively low
value compared to LPG and fuel gas.
272
-------
A. Cyclic Regeneration Scheme
FLUE GAS •«*
U>
OVERM.AI)
f. . \
STABiLiZER
HEFORMATE
Figure F4-18. Process flow diagram - catalytic reforming.
Source: Reference 7
-------
B. Semi-Regeneration Scheme
N5
-vl
.p-
COMPRESSOR
HEATERS AND REACTORS
/\ A
H2-RICH
STREAM
NET HYDROGEN
TO REFINERY
HYDROGEN TO
NAPHTHA HYDROTREATER
PRODUCT
H I SEPARATOR
TO FUEL GAS
^TO LIGHT-ENOS
RECOVERY
DESULFURIZED
NAPHTHA
C5 REFORMATS-»-
Figure F4-18. Continued.
Source: Reference 7, p. 165
-------
to
-^i
Ul
C. Continuous Regeneration Scheme
REGENERATOR
REACTORS
HO. 1
NO. 2
T_ ^ NO. 3
REGENERATED
CATALYST
SPENT CATALYST
HEATER LOW-PRESSURE SEPARATOR HIGH-PRESSURE SEPARATOR
COMPRESSOR
DESULFURIZEO
NAPHTHA CHARGE
H2-RICH
GAS
OVERHEAD
STABILIZER
REFORMATS
70-1525-?
Source: Reference 7
Figure F4-18. Continued
-------
^- • Small unit size or limited plot__S£ace.
• Unit shutdown for catalyst regeneration
at periodic intervals is acceptable.
• High quality feeds are available.
Factors favoring cyclic operation include:
• Octane numbers greater than 98 RON clear
are desired.
• Larger capacity units.
• Gasoline and hydrogen have high values
compared to LPG and fuel gas.
• Poor quality feeds are available.
• Return on investment more important than
initial investment.
• Greater operating flexibility.
The introduction of bimetallic catalysts (1968) al-
lowed reformers to be operated at lower pressures (higher sever-
ity) , particularly semi-regenerative units. Conventional
platinum catalysts, which dominated the field through the
1960's, are quite sensitive to coke deposits. Bimetallic
catalysts retain their activity and selectivity-at higher coke
levels and permit operation at lower pressures and lower hydro-
gen recycle ratios without sacrificing cycle length. The most
common "bimetallie combination is platinum and rhenium although
other combinations are available.5'6
276
-------
^- The trend toward even higher severity^lower pressure
operation led to the development of continuous reforming. A
continuous reformer utilizing three reactors is diagrammed in
Figure F4-18. Catalyst is continuously withdrawn (in small
batches) from the bottom of the third reactor while the unit is
in operation. The catalyst is regenerated in a separate regen-
eration system and returned to the top of the first reactor.
In four reactor systems, the fourth reactor stands
beside the stack containing the first three reactors. Catalyst
is withdrawn from the bottom of the third and fourth reactors,
sent to the regeneration system, and returned to the top of the
first and fourth reactors. In four reactor systems, the first
three reactors and the fourth reactor contain approximately
equal amounts of catalyst.8
Desirable features of continuous reforming include:8
• Both the reactor and regenerator systems
operate continuously and the unit need not
be shut down for periodic catalyst
regeneration.
• High severity-low pressure operation for
maximum product octane is possible since the
catalyst activity and selectivity is
continuously restored.
• A continuous stream of high purity hydrogen
is available to other process units.
277
-------
• The catalyst is not regenerated-within the
reactor; hence, both the reactor and re-
generator systems can be designed and operated
to perform their functions under optimum conditions
Economical operation at low pressures necessitated
some changes in the process flow and mechanical design of the
continuous reformer. As indicated in Figure F4-18, the cooled
reaction products flow into a low pressure separator where the
gas and liquid phases are separated. These two streams are
remixed at higher pressure, cooled, and separated in a high pres-
sure separator. Hence, high pressure separation is used in
a low pressure process. This increases the hydrogen purity and
reduces the amount of light hydrocarbons recycled to the re-
actors. In addition, the pressure drop across the recycle
hydrogen loop was minimized to reduce recycle compressor costs.8
One additional feature of the continuous reforming
process is that it can be installed without the regeneration
facilities as a semi-regenerative process. This substantially
reduces initial investment costs. Then, when higher octane
products are required, the unit can be converted to continuous
operation by installation of regeneration facilities.
Since the initial development of catalytic reforming,
many proprietary processes have been offered to the industry.
Many of these are available as either cyclic or semi-regenera-
tive processes. However, only UOP and IFF offer continuous
processes at present.8'9
A breakdown of current reforming capacity by the
method of regeneration is given in Table F4-45.-
278
-------
TABLE F4-45. REGENERATION SCHEMES USED. E.OR U.S.
~ REFORMING CAPACITY - ^- -
Regeneration Method
Semi-regeneration
Conventional Catalyst
Bimetallic Catalyst
Cyclic Regeneration
Conventional Catalyst
Bimetallic Catalyst
Other1
Conventional Catalyst
Bimetallic Catalyst
Number of
Units
42
125
21
7
3
3
201
Total Capacity,
bbl/sd
522,200
2,081,445
687,100
236,300
63,150
93,000
3,683,200
:0ther includes nonregenerative systems, i.e., catalyst replaced by fresh
catalyst, and continuous reforming processes.
Source: Reference 3.
Processing Variables--
Feedstock - Catalytic reformer feedstocks are usually
saturated (i.e., not olefinic) materials boiling up
to a maximum of 375°F. Generally, C5 and lighter
materials are not included in catalytic reformer
feeds as only the isomerization reactions lead to
octane improvement. Equilibrium concentrations of
the isoparaffins are quite low at process tempera-
tures.
279
-------
The most common feedstock is straigUt^run naphtha.
Other feedstocks include hydrocracked~"haphtha or
suitably treated coker naphtha.
The yield of gasoline of a given octane number at
given operating conditions depends on the type of
hydrocarbons present in the feed. Highly naphthenic
feeds, which convert readily to aromatics, are the
easiest to reform. Paraffin stocks, however, are more
difficult to reform since they depend on the more dif-
ficult isomerization, dehydrocyclization, and hydro-
cracking reactions. Hence, paraffin stocks give lower
gasoline yields than naphthenic stocks.
Contaminants in the feed can poison platinum catalysts,
reducing gasoline yield and catalyst life. Some of
the more important contaminants include sulfur, nitro-
gen, metals, water, and chlorine.1
Hydrogen sulfide is a reversible poison for platinum
catalysts, particularly the bimetallics, and causes a
decrease in the dehydrogenation and dehydrocyclization
activity. Hydrogen sulfide is formed under reactor
conditions from sulfur compounds in the feed and can
accumulate in the recycle hydrogen stream. Therefore,
hydrodesulfurization of the reformer feed to protect
catalyst activity is quite common.
Nitrogen compounds in the feed are converted to am-
monia under reactor conditions. The ammonia neutral-
izes the acid sites on the catalyst causing a decrease
in isomerization, hydrocracking, and ^dehydrocycliza-
tion activity. Organic nitrogen can be a problem with
280
-------
such feeds as coker naphtha, and high_pressure hydro-
treating may be required for nitrogen~~removal.
Extremely low concentrations of metals such as arsenic,
lead, and copper can deactivate platinum catalyst.
These metals, however, can be removed when desulfuriz-
ing the feed as the metals are deposited on the de-
sulfurizer catalyst, which is less sensitive than the
platinum catalyst.
The immediate effect of an increase in either the
water or chlorine content of the feed is to increase
the hydrocracking activity of the catalyst. Thus, the
yield of gasoline is reduced due to the formation of
Ci-Ci» cracking products. Ci-Cif products are present
as diluents in the recycle gas stream. This reduces
the hydrogen partial pressure resulting in increased
coke formation on the catalyst. Some water content,
however, is necessary as excessive demethylation of
substituted aromatics has been reported with water
content in the feed of less than 10 ppm.l
Some reforming catalysts contain chlorine as the acid
function promoter. Excess water, under these condi-
tions, serves to strip chlorine off the catalyst. The
effect of this is a net reduction in hydrocracking
activity as opposed to an increase in the short term.-
With this type of catalyst, a balance between the
chloride content and water is maintained by adjusting
the water content and injecting organic chlorides into
the reactor train.l
281
-------
Temperature - Increasing the reactor., temperatures can
increase product octane by increasing~£he reaction
rates. Typical reactor temperatures range, from 850-
1000°F. At temperatures below 850°F, the reactions
are too slow, while at temperatures greater than
1000°F, hydrocracking reactions become excessive,
reducing gasoline yield. In addition, some thermal
cracking can occur above this temperature, leading to
coke formation on the catalyst.
Space Velocity - The space velocity (vol. feed per
hour per vol. catalyst) required to achieve a given
severity depends on the hydrocarbon types present in
the feed; higher space velocity for naphthenic feeds
and lower space velocity for paraffinic feeds. Gen-
erally, space velocities in the range of 1.5-3.0
v/hr/v are used for gasoline production.*
Pressure - As mentioned previously, a reduction in the
reactor pressure increases the yield of high octane
aromatics by increasing the aromatic equilibrium con-
centration. However, coke formation is increased when
pressure is reduced since less hydrogen is available
to suppress coke formation.
Since the development of the catalytic reforming
process, there has been a trend to reduce operating
pressures to obtain substantial yield and octane bene-
fits. The early units, utilizing platinum catalysts,
showed rapid coke deactivation. These processes were
forced to operate at pressures in excess of 500 psig
to provide acceptable run lengths. The introduction
of bimetallic catalysts resulted in aMecrease in the
allowable reactor pressure. These catalysts appear
282
-------
to be more effective in dispersing the coke formed
during the reaction and retain their irctivity at much
higher coke levels. Thus, semi-regenerative units
utilizing bimetallics can retain acceptable run
lengths at pressures ranging from 150^200 psig.
Cyclic and continuous regenerative units now operate
at even lower pressures, ranging from 90-150 psig, due
to more frequent regeneration.
Hydrogen/Oil Ratio - The hydrogen to oil ratio changes
the hydrogen partial pressure in the reactor, and its
effects are therefore related to total pressure.
Typical hydrogen/oil ratios range from 2.0-5.5.
A summary of typical operating conditions along with
utility information is given in Table F4-46.
TABLE F4-46. OPERATING CONDITIONS FOR CATALYTIC REFORMING
Reactor Pressure, psig
Semi-regenerative 150-500
Cyclic Regenerative 90-200
Continuous Regeneration 90-200
Reactor Temperature, °F 850-1000
Space Velocity, vol/hr/vol 1.5-3.0
Hydrogen to Hydrocarbon Ratio 2.0-5.5
Hydrogen Recycle Purity, vol. % 80-85
Utilities - per barrel feed
Power, kwh 5-7
Fuel, 106 Btu 0.15-0.32
Source: -_ References 1, 2,4
283
-------
4.5.1.3^- Process Emissions-- _. ^
~ -f- -
Emission sources from catalytic reformers .include:
• Catalyst regeneration operations,
• Process heater flue gas, and
• Fugitive emissions.
Catalyst Regeneration--During the reforming operation,
coke is deposited on the catalyst. The rate of coke formation
is a function of the type of feedstock and the severity of the
operating conditions. This coke must be removed from the cat-
alyst to restore catalytic activity.
The regeneration procedure consists of burning these
deposits off the catalyst by injecting air into the reactor.
During this process, a flue gas stream is generated which con-
tains carbon monoxide and low concentrations of sulfur and
nitrogen oxides. Total mass emissions from regeneration, how-
ever, are quite low because only small amounts of coke are
produced, and the frequency of regeneration may be low.
Perhaps the highest potential for emissions from
regeneration operations occurs with the so-called continuous
reformers, since they can be operated at conditions of highest
severity. Carbon monoxide emissions from continuous reformers
have been estimated at 0.002-0.02 pounds per barrel of fresh
feed.12 Hence, emissions from this source are relatively in-
significant.
284
-------
^- Process Heaters--Emission factors fgr_various pollu-
tants from fired heaters are given in Table F4~-37. A typical
reforming unit will require a process heater to heat the feed
stream prior to entering each of the reactors. In some cases,
one large heater is used for this purpose. In addition, a
heater may be required as a stabilizer reboiler. Heat input re-
quirements vary, ranging from 150,000-320,000 Btu per barrel
feed.
Total emissions from the reformer heaters are also
given in Table F4-47. These figures are given as pounds of
pollutant per thousand barrels of fresh feed using a heat input
of 200,000 Btu per barrel feed.
Fugitive Emissions — Fugitive emissions can occur from
a variety of sources including valves, pumps, compressors,
flanges and other fittings, relief valves, and drains. Hydro-
carbon emission rates for each individual source are quite low.
However, total fugitive emissions are significant because of
the large number of such sources within the refinery.
Total estimated fugitive non-methane hydrocarbon emis-
sions from a typical catalytic reforming unit are given in
Table F4-48. The listed emission factors were determined as a
part of this program. Additional information on these emission
factors, including a complete discussion on their derivation
and the confidence intervals for each source category, are con-
tained in Appendices B and C of this report.
Estimates for the number of sources within each source
category were developed from:
285
-------
TABLE Pl-47.
TYPICAL EMISSIONS FROM CATALYTIC REFORMING
UNIT PROCESS HEATERS
EPA Emission Factor3
(lb/103 gal-oil fired)
(lb/106 scf-gas fired)
Total Emissions
(lb/103 bbl of
fresh"feed)
Oil Fired Heaters
Particulates
- Distillate oil
- Residual oil
Grade 4
Grade 5
Grade 6
Sulfur Dioxide
- Distillate oil
- Residual oil
Sulfur Trioxidec
Carbon Monoxide
Hydrocarbons (as CHi,)
Nitrogen Oxides
(as N02)
- Distillate oil
- Residual oil6
7
10
10(S)+3
142(S)
157(S)
2(S)
5
1
22 2
22+400 (N)
2.9
10
14
14(S)+4.3
203(S)
224(S)
2.9(S)
7.1
1.4
31
Gas Fired Heaters
Particulates
Sulfur Oxides (as S02)£
Carbon Monoxide
Hydrocarbons (as CHi*)
Nitrogen Oxides (as K02)
5-15
0.6
17
3
120-230
0.95-2.9
0.11
3.2
0.57
22.9-43.8
Source: Reference 11
Eased on a heat input of 200,000 Btu/bbl of fresh feed with the following
fuel heating values: Oil - 140,000 Btu/gal; Gas - 1050 Btu/scf.
CS • wt Z sulfur in the oil
laproper combustion may cause a significant increase in emissions
Uae this emission factor for residual oils with less than 0.5% Qf<.5) nitro-
gen content. For oil with higher nitrogen content QN>0,5}, use emission
factor of 120 lb/103 gal
Based_on sulfur content of 2000 gr/10G scf —
286
-------
TABLE F4-48.
ESTIMATED FUGITIVE NON-METHANE HYDROCARBON EMISSIONS FROM
A TYPICAL CATALYTIC REFORMING UNIT
to
00
--J
Emissions
Source Type
Valves
Open-End
(Sample)
Valves
Pumps (Pump
Seals)
Drains
Flanges &
Fittings
Relief Valves
Compressors
(Compressor
Sefnls) 1
Process
Stream Service
Classification
Gas /Vapor
Light Liquid
(VP > 0.1 psla ? 100'F)
Heavy Liquid
(VP $ 0.1 piila 9 100'F)
Hydrogen Service
Total
All
Light Liquid
(VP > 0.1 psla 1? 100'F)
Heavy Liquid
(VP < O.I pala 9 100'P)
Total
All
All
All
Hydrocarbon
Hydrogen,
Total '
Number of Sources in
Counts or Estimates
From Radian Study
IfiO
391
43
77
691a
_
13(18)
K 2)
14(20)a
49a
296la
6C
0(0)
3(6)*
Process Unit
Counts or Estimates
From PES Study
154 - 291
493 - 938
0
139 - 263
786 - 1492C
16 - 30*>
8(11)-17(24)
0
8(ll)-17(24)b
-
2585 - 4935C
• -
0(0)
3(6)
3(6)
Source
Emission
Factor, Ib/hr
0.059
0.024
0.0005
0.018
0.005
0.25
0.046
0.070
0.00056
0.19
1.4
0.11
Estimated Total
Emissions,
Ib/hr
9.09 - 17.2
9.38 - 22.5
0.0
1.39 - 4.73
19.9 - A4.4
0.080 - 0.15
2.75 - 6.00
0.0 - 0.092
2.75 - 6.09
3.43
1.45 - 2.76
1.14
0.0
0.66 -
i ' Of66
29.4 - 58.6
Physically Counted
Counted Fron Flow Diagrams
CEstlmated
Reference 10
-------
. • Actual field source counts conducted during
the course of this program, and ~~ "
• Counts contained in The Development of
Petroleum Refinery Plot Plans, EPA pub-
lication number 450/3-78-02510
The methods used to develop estimates for each source
type have been previously described in Section 4.1.1.3 of this
appendix.
The source counts described in Table F4-48 refer to
those sources located within the battery limits of the process.
That is, equipment located in tankage or transfer lines is not
included in these source counts.
The number of sources per unit depends on a variety
of factors including processing complexity, process type, and
processing flexibility. Hence, these source counts may not be
representative of all or even the majority of reforming units.
The estimated composition of fugitive non-methane
hydrocarbon emissions from a catalytic reforming unit is given
in Table F4-49. These estimates indicate the types of hydro-
carbons emitted by sources on a number of process streams found
within the unit. Additional information about these estimates
may be jfpund in Section 2.4 of Appendix D.
4.5.2 Alkylation
Alkylation refers to the chemical combination of two
hydrocarbon molecules, an olefin and an isopara'ff in, to produce
higher--molecular weight isoparaff ins . Typical dlefin feedstocks
288
-------
TABLE'F4-49.
hO
oo
ESTIMATED COMPOSITION OF FUGITIVE NON-METHANE
EMISSIONS FROM A CATALYTIC REFORMING UNIT
Stream
Estimated percentage of emissions
attributed to each stream - Z
Weighted contribution of each
component to unit emissions - ppmw
Benzene
Toluene
Ethylbenzene
Xylenes
Other Alkylbenzenes
Naphthalene
Anthracene
Biphenyl
Other Polynuclear aromatlca
n-Hexane
( ' Othar Alkanea '
Olefins
Cyclo Alkanes
Hydrogen
Desulfurlzed
Naphtha
47
119
1232
417
763
7792
688
2
295
7042
18254
234817
0
198579
0
Re formate
47
2538
36519
15745
80323
152468
3478
0
0
329
11280
167320
1
0
0
0
H2 Recycle
Gas
6
0
0 .
0
0
0
0
0
0
0
0
39000
0
0
21000
Totals
1001
2657
37751
16162
81086
160260
4166
2
295
7371
29534
441137
i
0
198579
21000
1000000
-------
consist^of butylene or a mixture of propylene^and butylene,
.aT • -
while the isoparaffin most commonly used is isoButane. The
resulting product is a branched isoparaffin containing from six
to nine carbon atoms which has a high octane rating. The
alkylate product is usually used to upgrade the octane rating
of the motor gasoline pool.
The growth of alkylation to its present status as a
major refining process is due to the excellent characteristics
of the alkylate product. It has a high octane rating, a high
heat of combustion, a low vapor pressure, and a desirable
boiling range. In addition, the alkylate has good lead suscep-
tibility with low or even negative sensitivity (RON-MON) . l "*
The first commercial alkylation unit went onstream in
1939 and by 1946, 59 alkylation units were in operation. These
units helped to provide the large quantities of aviation fuel
required by the military during World War II. After the war,
many of these units were shut down or dismantled because the
octane levels required for motor gasoline were not high enough
to justify the continued use of this relatively expensive
process.:4
Alkylation capacity began another sharp increase dur-
ing the Korean War. With ever increasing octane requirements
for motor gasoline, alkylation capacity continued to expand.
Today, -the total U.S. alkylation capacity is in excess of
900,000 barrels per day.3
4.5.2.1 Process Description--
Almost all of the commercial alkylation units utilize
either~sulfuric acid (H2SOi,) or hydrofluoric acid (HF) as
290
-------
catalyst?.- The sulfuric acid process was offered to the indus-
-j£- • -
try in the late 1930's as a means of utilizing light olefins
produced in the thermal and catalytic cracking processes. This
early experimental work indicated that the alkytation reactions
would proceed in the absence of a catalyst at elevated tempera-
tures and that catalysts other than sulfuric acid (such as
hydrofluoric acid or aluminum chloride) could also be used.
Successful processes were developed using hydrofluoric acid
catalyst and the first commercial unit was built in 1942. The
thermal and aluminum chloride processes, however, found only
limited commercial acceptance. * **
The current use of sulfuric and hydrofluoric acid
alkylation processes by U.S. refiners is given below in Table
F4-50.
TABLE F4-50. U.S. ALKYLATION CAPACITY, 1979
Process
Sulfuric Acid Alkylation
Hydrofluoric Acid Alkylation
Number
of Units
61
66
127
Capacity,
bbl/sd
522,100
384,100
906,200
Source: Reference 3
4.5.2.2__ Process Technology
Sulfuric Acid Alkylation--Process flow diagrams for
two typical sulfuric acid alkylation units are given in Figures
F4-19 and F4-20. The olefin feed, recycle and makeup isobutane,
and acid are mixed within a specially designed Teactor. In the
reactor"; the olefin and isobutane combine to form higher
291
-------
SATURATE BUTANES I
to
BUTANE
FEED
OLEFINS
DtPROPANIZER OEISOBUTANIZER
TREATER
EFFLUfNT
TRFATER
FRESH AC 102
I
SPENT
RVP AUYLATE
PROPANF
70-1022-1
Figure F4-19. Process flow diagram - H2SO., aLkylation,
cascade auto-refrigeration system.
Source: Reference 18, p. 69
-------
NO
VD
I inmf Miir \
nu i m nip ? ^J.
Figure F4-20. Process flow diagram - H2SCK alkylation,
Stratco effluent refrigeration system.
Source: Reference 18, p. 68
-------
molecular- weight isoparaff ins. The contents of J:he reactor
*<" -
are thoroughly agitated to promote contact betwe~&n -the hydro-
carbon phase and the acid catalyst. The emulsion is.then sent
to an acid settler where the hydrocarbon and acid form separate
phases. The hydrocarbon phase is caustic treated and/or water
washed and charged to a deisobutanizer for fractionation into
an isobutane recycle stream, a normal butane stream, and the
alkylate product. Part of the acid phase is returned to the
reactor while the remainder is removed from the system as spent
acid. In some cases, this acid is regenerated and returned to
the system.
Reactor design - The sulfuric acid alkylation reactor
must be designed to perform several important func-
tions including:
• Removal of heat generated by the re-
action, and
* Maintenance of a high isobutane to
olefin ratio.
The heat of reaction ranges from 270-360 Btu per
pound of alkylate. This heat must be removed from
the reactor to maintain the desired reactor tempera-
ture. High isobutane to olefin ratios are important
in order to minimize side reactions such as the poly-
merization of the olefins.15
Several types of reactors have been d&signed to meet
these requirements. And, the two most prevalent
designs are the Kellogg auto-refrigera_ted or cascade
294
-------
reactor and the Stratco stirred cont_aetor. These
are shown in Figures F4-19 and F4-26". ~~" "
In the cascade system, a dilute olefin concentration
is obtained by splitting the olefin feed into a num-
ber of parallel streams. Each stream is fed to a
separate reactor .compartment, each with its own mixer.
The isobutane-acid mixture flows in series from one
reactor compartment to the next and ends up in the
acid settler. This flow scheme provides the highest
effective isobutane concentration for the entire re-
actor.
Low reactor temperatures are maintained by allowing a
portion of the reactor contents to vaporize, thus re-
moving heat from the reaction mixture. The vapors
are compressed and charged to a depropanizer to pre-
vent propane from accumulating within the system.
Butanes from the depropanizer bottoms are combined
with recycle isobutane from the deisobutanizer and
returned to the reactor.lk
The Stratco system utilizes external cooling to re-
move heat generated during the reaction. The reactor
contains a tube bundle similar to that of a heat ex-
changer. The reactor contents are mixed and circu-
lated around the tubes by an internal- axial flow
impeller.l 5
The hydrocarbon phase from the acid settler is passed
through the tube bundle within the reactor. Much of
the isobutane is allowed to flash, dropping the tem-
perature of this stream some 20-30°F.^ And, heat
295
-------
generated within the reactor is removed by heat ex-
change with this stream. The isobutane vapors are
compressed, depropanized, and returned to the reactor
along with isobutane from the deisobutanizer.3'*u
Both of the above reactor systems maintain high iso-
butane to olefin ratios to prevent side reactions. In
the cascade system, the olefin feed is split into
several streams. Each stream is introduced into a
separator reactor compartment which keeps the olefin
concentration low during the entire reaction sequence.
In the Stratco system, the olefin feed is rapidly
dispersed by the mixer. None of the isobutane in the
reactor need be vaporized for cooling. Hence, the
isobutane concentration remains high.
Process variables - The most important process vari-
ables in sulfuric acid alkylation are reaction tem-
perature, acid strength, isobutane concentration, and
olefin space velocity.2
The reaction temperature is usually kept between 35-
55°F. At lower temperatures, sulfuric acid becomes
so viscous that good mixing in the reactor and sub-
sequent phase separation of the emulsion become very
difficult. At higher temperatures, polymerization of
the olef ins causes yields to decrease-. 2
The highest quality alkylate is obtained using acid
strengths of 93-95 percent by weight.- A decrease in
acid strength lessens its catalytic activity and a
decrease in product octane results.
296
-------
^- The isobutane concentration is usually^, expressed as
the isobutane to olefin ratio. High"~isobutane to
olefin ratios increase the octane of the alkylate,
minimize the number of side reactions, and decrease
the Level of acid consumption. The external isobutane
to olefin ratio generally ranges from 5:1 to 15:1.
However, internal ratios may reach as high as 1000:1
in well mixed reactors.2 Units with higher external
isobutane to olefin ratios require more investment
and consume more utilities. These costs are associ-
ated with the increased stream flows requiring larger
compressors, lines, and fractionation systems.
In general, the lower the olefin space velocity, the
higher the octane of the alkylate. Lower space velo-
city can also reduce acid consumption and decrease the
amount of high boiling hydrocarbons produced by side
reactions. Space velocities for sulfuric acid alkyl-
ation range from 0.1-0.6 v/hr/v.2 A summary of
sulfuric acid alkylation process conditions along
with utility requirements are given in Table F4-51.
Hydrofluoric Acid Alkylation--A flow diagram for a
typical hydrofluoric alkylation unit is given in Figure F4-21.
Olefin feed (Cs-Ct,) and isobutane are mixed with acid within a
specially designed reactor. The resulting emulsion is sent to
an acid_settler for separation of the hydrocarbon and acid
phases. As is the case with sulfuric acid alkylation, the heat
produced by the reaction must be removed from the reactor to
maintain the desired temperature. And, the method of heat re-
moval depends on the reactor design.17
297
-------
MIXER-SETTLER ISOSTKll'PER
ACID REGENERATOR
DEPROPANIZER HF STRIPPIK
OO
REAC1OR s
OLEFIN
FEED
KOH TRCATERS
70-1521-1
Figure F4-21. Process flow diagram - HF alkylation.
Source: Reference 18, p. 67.
-------
TABLE F4-51. OPERATING INFORMATION^FOR
-SULFURIC ACID ALKYLATlON
Isobutane Concentration
Vol. % in reaction zone 40 - 80
External ratio to olefins 3-12:1
Internal ratio to olefins 50-1000:1
Olefin Concentration
Total hydrocarbon contact time, min. 20 - 30
Olefin space velocity, v/hr/v 0.1 - 0.6
Reactor Temperature, °F 35 - 60
Sulfuric Acid Concentration, wt % 88 - 95
Acid in Emulsion, wt % 40 - 60
Reactor Pressure
Stratco contactor system, psig >50
Kellogg cascade system, psig 5-15
Utility Data - per bbl total alkylate
Steam, Ib 300 - 400
Power, kwh 2.5 - 5.0
Chemical Requirements - per bbl total alkylate
HzSOL,, Ib 18 - 30
Caustic, Ib 0.2
Source: References 1, 2, 4, 18
The hydrocarbon phase in the settler contains the
alkylate product, excess isobutane, and a small amount of acid.
This mixture is sent to an isostripper for fractionation into
a recycle isobutane stream and the alkylate product. As indi-
cated in Figure F4-21, field butanes (a mixture of butane
isomersj can also be charged to the isostripper. Hence, outputs
from the isostripper include an isobutane overhead stream, a
normal butane stream, and the alkylate bottoms product.16
299
-------
The normal butane stream is withdrawn from the iso-
stripper as a vapor sidecut. In addition to removing, normal
butane from the system, removing the butane in this manner pro-
vides control over the vapor pressure of the alkylate product.16
A portion of the isostripper overhead is sent to a
depropanizer to prevent accumulation of propane within the sys-
tem. The depropanizer overhead contains a small amount of acid.
This acid is removed in an HF stripper and returned to the re-
actor. J 6
Most of the acid phase in the acid settler is recycled
back to the reactor. However, a slipstream of acid from the
settler is sent to an acid regenerator to effect separation of
the acid from water which enters with the feed and polymers
which are formed in small volumes in the reactor. The column
produces reconstituted acid as an overhead product which is
returned to the reactor. The bottoms product, containing the
polymer and water, is commonly burned as fuel.2
Corrosion in HF alkylation is largely iron fluoride
deposition and stress corrosion. Hydrogen fluoride attacks the
slag formations in welds, and process vessels must be stress
relieved. Any part of the plant which is in direct contact with
HF is made of monel metal.1 An additional corrosion prevention
method is the dehydration of the olefin and isobutane feeds over
a solid hed desiccant. This operation reduces the water in the
feed to less than one percent.
Defluorination of the alkylate product^can be accom-
plished by proper design of the isostripper reboiler. The
alkylate is passed through the tubes of a fired-heater and the
concentration of combined fluorides decreases, often to less
than ICTppm.
300
-------
Reactor design - Most of the HF units, gresently in
operation have been designed by either"TJOP Inc. or
Phillips Petroleum Co. The reactors designed by UOP
are similar to shell and tube heat exchangers. Cool-
ing water flows inside the tubes to remove the heat
of reaction. Mixing of the emulsion is accomplished
by using an external recirculation pump to force the
mixture through the reactor.15
The reactors designed by Phillips can be described
as vertical lift, plug flow reactors. Hydrocarbon
feed enters the bottom of the vertical tube reactor
through a jet-eductor device and is contacted with
acid from an acid cooler. The jet-eductor helps to
mix the hydrocarbon and acid phases. The mixture
flows upward to a settler for separation of the acid
and hydrocarbon phases. The settled acid flows down-
ward to the acid cooler and is remixed with fresh
hycrocarbon. Acid circulation in this system is pro-
moted by the hydrocarbon jet-eductor and by the grav-
ity differential between the HF acid recycle leg and
the acid-hydrocarbon mixture in the reactor tube. In
this design, an acid recirculation pump is not neces-
sary.2' l 5
Process variables - The important variables in hydro-
fluoric acid alkylation include the isobutane concen-
tration, the reactor temperature, and the acid
strength. High isobutane concentrations are required
to produce high octane alkylate while minimizing side
reactions.
301
-------
reactor temperatures used in HF alkylation are
"generally between 70-100°F. The effects~of tempera-
ture, however, are not as critical as those encoun-
tered in sulfuric acid alkylation.2
In hydrofluoric acid alkylation, the highest octane
alkylate is obtained when using acid concentrations
of 86-90 percent by weight. In commercial operations,
acid concentrations range from 83-92 percent with less
than one percent water.
A summary of hydrofluoric acid alkylation operating
conditions and utility requirements given in Table
F4-52.
TABLE F4-52. OPERATING INFORMATION FOR HYDROFLUORIC
ACID ALKYLATION
Isobutane Concentration
Vol % in reaction zone 30-80
External ratio to olefins 3-12:1
Hydrocarbon Contact Time, min 8-20
Reactor Temperature, °F 60 - 120
Hydrofluoric Acid Concentration, wt °L 83 - 92
Acid in Emulsion, wt % 25-80
Utility Data - per barrel total alkylate
Power, kwh 3-7
Fuel, 106 Btu 0.3 - 1.1
Chemical Data - per barrel total alkylate
Acid, Ib 0.1 - 0.5
Caustic, Ib 0.1-0.2
Source^ References 1, 2, 4, 18
302
-------
4.5.2.3 . ^Process Emissions--
The alkylation process is a closed system; that is,
there are no process vents to the atmosphere. The only air
emissions are those associated with process heaters and fugitive
emissions from process equipment and fittings.
Process Heaters--Emission factors for various pollu-
tants from fired heaters are given in Table F4-53. Typically,
an HF alkylation unit will require a process heater. It is used
as a reboiler for the main fractionator. Heat input require-
ments vary, ranging from 300,000 to 1,100,000 Btu per barrel of
total alkylate."»18
Total emissions from the alkylation heater are also
given in Table F4-53. These figures are given as pounds of
pollutant per thousand barrels of total alkylate using a heat
input of 360,000 Btu per barrel total alkylate.12
Fugitive Emissions—Fugitive emissions can occur from
a variety of sources including valves, pump seals, compressor
seals, flanges and other fittings, relief valves, and drains.
Hydrocarbon emission rates for each individual source are quite
low. However, total fugitive emissions are significant be-
cause of the large number of such sources within the refinery.
-.- Total estimated fugitive non-methane hydrocarbon emis-
sions from a typical sulfuric acid alkylation unit are given in
Table F4-54. The listed emission factors were determined as a
part of this program. Additional information on_these emission
factors, including a complete discussion on their derivation and
the confidence intervals for each source category, are contained
in Appendices B and C of this report.
3Q3
-------
TABLE F4-53.
TYPICAL EMISSIONS FROM ALKYLATION
UNIT PROCESS HEATERS '~ ~
EPA Emission Factor
(lb/103 gal-oil fired)
(lb/106 scf-gas fired)
Total Emissions
(lb/103 bbl of
total alkylate)
Oil Fired Heaters
Particulates
- Distillate oil
- Residual oil
Grade 4
Grade 5
Grade 6
Sulfur Dioxide0
- Distillate oil
- Residual oil
Sulfur Trioxidec
Carbon Monoxide
Hydrocarbons (as
Nitrogen Oxides
(as N02)
- Distillate oil
- Residual oile
7
10
10(S)+3
142(S)
157(S)
2(S)
5
1
22
22+400(N)
5.1
18
26
26(S)+7.7
365 (S)
404 (S)
13
2'.6
57
57+1030(N)'
Gas Fired Heaters
Particulates
Sulfur Oxides (as S02)f
Carbon Monoxide
Hydrocarbons (as CHi»)
Nitrogen Oxides (as N02)
5-15
0.6
17
3
120-230
1.7-5.1
0.21
5.8
1.0
41.1-78.9
Source: Reference 11
bBased on a heat input of 360,000 Btu/bbl of total alkylate with the following
fuel_h.eating values: Oil - 140,000 Btu/gal; Gas - 1050 Btu/scf.
CS • wt Z sulfur in the oil
Improper combustion may cause a significant increase in emissions
*Use this emission factor for residual oils with less than 0,5Z QK.5). nitro-
gen content. For oil with higher nitrogen content QJ>0,5}, use emission
factor of 120 lb/103 gal
Based-«n sulfur content of 2000 gr/106 scf
304
-------
TABLE FA-54. ESTIMATED FUGITIVE NON-METHANE HYDROCARBON EMISSIONS FROM
A TYPICAL SULFURIC ACID ALKYLATION UNIT
1 i ,"
Emissions
Source Type
Valves
Open-End
(Sample)
Valves
Pumns (Pump
Seals)
Drains
Flanges &
Fittings
Relief Valves
Compressors
(Compressor
,Seals) |
Process
Stream Service
Classification
Gas /Vapor
Light Liquid
(VP > 0.1 pala « 100 '?)
Heavy Liquid
(VP < 0.1 pala « 100'F)
Hydrogen Service
Total
All
Light Liquid
(VP > 0.1 pata 9 100'F)
Heavy Liquid
(VP < 0.1 p«la « 100'P)
Total
All
All
All
Hydrocarbon
Hydrogen
Total
Number of Sources In Process Unit
Counts or Estimates
From Radian Study
274
403
0
0
677a
_
13(18)
0( 0)
13(18)a
41*»
2407a
6C
0
0
0
Counts or Estimates
From PES Study
429 - 719
636 - 1067
0
0
1065 - 1786C
26 - 30b
13(18)-23(32)
0( 0)- 0(0)
13(18)-23(32)b
-
3525 - 5875C
-
0(0) - 2(4)
0(0) - 0(0)
0(0) - 2(4)b
Source Estimated Total
Emission Emissions,
Factor, Ib/hr Ib/hr
0.059
0.024
0.0005
0.018
0.005
0.25
0.046
0.070
0.00056
0.19
1.4
0.11
t
16.2 - 42.4
9.67 - 25.6
0.0
0.0
25.9 - 68.0
0.13 - 0.15
4.50 - 8.00
0.0
4.50 - 8.00
2.87
1.35 - 3.29
1.14
0.0 - 5.60
0.0
0.0 - 5.60
35.9 - 89.1 '
j . ii '
aPhyalcally Counted
Counted From Flow Diagrams
CEstiraoted
Reference 10
-------
f-stimates for the number of sources within each source
category were developed from: •- -^ .
• Actual field source counts conducted during
the course of this program, and
• Counts contained in The Development of
Petroleum Refinery Plot Plans, EPA pub-
lication number 450/3-78-025lc
The methods used to develop estimates for each source
type have been previously described in Section 4.1.1.3 of this
Appendix.
The source counts presented in Table F4-54 refer to
those sources located within the battery limits of the process.
That is, equipment located in tankage or transfer lines is not
included in these source counts.
The number of sources per unit depends on a variety of
factors including processing complexity, process type, and
processing flexibility. Thus, the number of sources presented
in Table F4-54 may not be representative of all or even the
majority of alkylation units.
The estimated composition of fugitive non-methane
hydrocarbon emissions from an alkylation unit is given in
Table F4-55. These estimates indicate the types of hydrocarbons
contributed by sources on a number of process lines within the
unit. Additional information about these estimates may be found
in Section 2.4 of Appendix D.
306
-------
TABLE F4-55
ESTIMATED COMPOSITION OF FUGITIVE NONMETHANE
HYDROCARBON EMISSIONS FROM AN ALKYLATION UNIT
LO
o
Estimated percentage of emissions
attributed to each stream - %
Weighted contribution of each
component to unit emissions - ppmw
Benzene
Toluene
Ethylbenzene
Xylenes
Other Alkylbenzenes
Naphthalene
Anthr.nr.ene
Blphenyl
Other I'olynuclear aromatics
' n-Hc'xane
Other Alkanes
Olefina
Cyclo Alkanes
LPG
Olefins
Ik
0
0
0
0
0
0
0
0
0
0
96000
144000
0
Stream
LPG
35
0
0
0
0
0
0
0
0
0
0
350000
0
0
H2SO^
0
0
0
0
0
0
0
0
0
0
0
0
0
0
Alkylate
41
0
0
0
1
1
0
0
0
1
39
40957'2
. 381
. 5
H.
Total s
0
0
0
1
1
0
0
0
1
/ > 39
8^72
144381
5
i(5ooooo
-------
4.5.3 Isomerization
• •r •_
Isomerization processes are used to convert normal
paraffins into isoparaffins. The value of this type of pro-
cessing is indicated in Table F4-56, which is a listing of the
octane number of various light hydrocarbons. In general, the
octane numbers of the isoparaffins are considerably higher than
those of the normal paraffins.
TABLE F4-56. OCTANE NUMBERS OF PARAFFIN HYDROCARBONS
Paraffin
Hydrocarbon
n-butane
Isobutane
n-pentane
Isopentane
n-hexane
3-methylpentane
2-methylpentane
2,3-dimethylbutane
2,2-dimethylbutane
RON
Clear
94.0
102.1
61.7
92.6
34.0
74.5
73.4
103.5
92.3
MOW
Clear
89.1
97.0
61.3
90.3
25.0
74.0
72.9
94.3
92.9
RON
3 cc TEL
104.1
118.3
84.7
103.5
65.3
92.3
92.2
112.0
103.4
MON
3 cc TEL
104.7
-
83.6
106.9
63.5
92.6
92.4
109.7
114.6
Source: References 1, 20
308
-------
4.5.3.1- ^'Process Description--
-*T ^
The isomerization process was originally developed and
used to provide isobutane for the alkylation process. Catalytic
cracking processes provided an abundant source of olefins, but
the amount of isobutane present in refinery streams was insuf-
ficient to meet alkylation feed requirements. Consequently,
isomerization was used to convert normal butane to isobutane.
Alkylation and isomerization processes helped to provide the
large quantities of high octane aviation fuel required during
World War II.20
The isomerization of pentane and hexane to their
isomers has become increasingly important for a number of rea-
sons. Until recently, refiners have had little reason to
isoiuerize Cs/Ct streams because the octane of these materials
could be significantly improved with the addition of lead com-
pounds. However, with the reduction in allowable lead content
and increasing octane requirements for modern gasolines, some
refiners have turned to isomerization as a means of utilizing
these low octane materials.21
The current domestic isomerization capacity, as a
function of the feedstock, is given in Table F4-57.
TABLE F4-57. CURRENT U.S. ISOMERIZATION CAPACITY (1979)
Feedstock
Butane
Pentane .
Pentane-4iexane
Number
of Units
19
5
3
Total Capacity,
bbl/sd
44,550
60,400
26,300
Source:" Reference 3
309
-------
4.5.3.2. ^-Process Technology-- _ ._
" -r- - ""-
Basically, the isomerization process consists of con-
tacting hydrocarbon and catalyst in a reactor, and then recover-
ing the products. The reactor conditions are se~t to give favor-
able reaction rates and equilibria with a minimum amount of side
reactions. Separation of the unreacted normal paraffins from
the isoparaffin product can be accomplished using distillation
or solid adsorbent systems.
The effect of temperature is very important in iso-
merization processes. In general, equilibrium concentrations
of the isoparaffins are increased by reducing the reaction temp-
perature.19'20
This effect is illustrated in Table F4-58 which shows
the equilibrium composition of several n-paraffin-isoparaffin
mixtures as a function of temperature.
The catalysts used for isomerization are of three
general types. The first group consists of aluminum chloride
promoted with hydrochloric acid. The reaction may be conducted
in the vapor phase over a solid catalyst bed, or in the liquid
phase with the catalyst dissolved in a solvent such as antimony
trichloride.
. ^-- The second group of catalysts contain -a metal such
as platinum on a silica-alumina or zeolite base. With these
catalysts, the reaction occurs in the vapor phase in packed-bed
reactors. These dual functional catalysts are nencorrosive and
31Q
-------
TABLE FZt-58. THERMODYNAMIC EQUILIBRIA OF BUTANE,
PENTANE, AND HEXANE MIXTURES - ^-
Compound,
Vol. %
n-butane
Isobutane
n-pentane
Isopentane
n-hexane
2-methylpentane
3-methylpentane
2 ,2-dimethylbutane
2,3-dimethylbutane
212°F
35
65
100
16
84
100
9
28
13
39
11
100
Temperature
570°F " 800°F
52
48
100
^38
^62
100
14 "-21
38 "-40
19 "-20
20 MO
9 - ^9
100 100
Source: Reference 1, 20
easy to handle. However, operating temperatures are higher
resulting in a less favorable equilibrium.19
The third group of catalysts are the low temperature,
dual-functional hydroisomerization catalysts. These catalysts
are essentially a hybrid of the first two groups, and are the
most recent types to be developed. These "third generation"
catalysts are prepared by treating a platinum-alumina composite
with a polyhalide such as carbon tetrachloride, chloroform,
aluminum chloride, etc. These catalysts permit operations in
the range of 200-400°F, temperatures which promote higher equi-
librium concentrations of the isoparaffin.*9'2°
311
-------
Figure F4-22 shows a flow diagram fof a typical solid-
bed butane isomerization unit. Butane feed is first charged to
a deisobutanizer where isobutane is removed as the overhead
product. The normal butane stream from the bottom of the deiso-
butanizer is mixed with hydrogen to suppress the polymerization
of small amounts of olefins formed during the isomerization
reaction. Hydrogen consumption, however, is quite low.2 The
mixture is sent to a fired heater, and the vaporized feed is
charged to the reactor. In the reactor, the n-butane is con-
verted to a near equilibrium mixture of normal and isobutane.
The conversion to isobutane is maximized by operating at low
temperatures. In addition, trace amounts of organic chlorides
are added to the reactor feed to promote the reaction. The re-
actor effluent is cooled and sent to a separator where recycle
hydrogen is recovered. The mixture is stabilized to remove
traces of light gases and sent to the deisobutanizer for separa-
tion of the isobutane product. Unconverted n-butane plus n-
butane from the fresh feed are returned to the system.
As mentioned previously, butane isomerization is a
means of providing isobutane feed for alkylation. These two
units can be designed to use common distillation equipment
with a considerable reduction in cost.
The design of C5-C6 fixed bed reactor isomerization
units is similar to that of Ci, isomerization units; however,
several' differences are noted. First, the feed is usually
hydrotreated and dried with molecular sieves to protect catalyst
activity. In addition, a guard or lead isomerization reactor
may be. used to saturate aromatics and olefins present in the
feed.
312
-------
OEISOBUTANIZER
STABILIZER
REACTOR
OJ
I—1
U)
BUTANE
FKtn
HYbROGEN i
*• CAS TO rUEL
*• ISOBUTANE
COMPRESSOR
70-1510-1
Figure F4-22. Process flow diagram - butane isomerization.
Source: Reference 18, p. 66.
-------
. i,- A flow diagram for the n-butane liquid phase isomeri-
.j£- '
zation process is given in Figure F4-23. This "pro'cess is based
on the use of liquid aluminum chloride, promoted with hydrogen
chloride, as the catalyst. The n-butane feed i"s passed through
a feed drier, a heater, and a catalyst scrubber tower to the
reactor. A small sidestream of catalyst from the reactor is
also charged to the catalyst scrubber column. The active
catalyst components are extracted with the fresh feed and re-
turned to the reactor. A fluid inactive aluminum chloride-
hydrocarbon complex, which forms as a result of side reactions,
flows to the bottom of this column and is drained off. Hydrogen
chloride recycle and makeup streams are added to the butane
feed. This mixture is sent to the reactor where it is contacted
with the aluminum chloride catalyst, which may be in the form of
a complex or dissolved in a solvent such as molten antimony
trichloride. After the hydrocarbon and the catalyst phases
separate in the upper section of the reactor, the hydrocarbon
phase is sent to the catalyst removal column for removal of the
dissolved catalyst by distillation. The recovered catalyst is
sent back to the reactor while the hydrocarbons are sent to an
HC1 stripper where acid gas is removed and recycled to the reac-
tor. The mixed butanes from the HCl stripper are caustic washed
and separated into normal and isobutane product streams.J
The flow plan for the liquid phase isomerization of
Cs-Ce fractions is similar to that for d, isomerization. How-
ever, a-hydrogen system is added to suppress disproportioning
reactions.
Operating conditions and utility information for the
isomerization process are given in Table F4-59.
314
-------
niiVERS
rttn
ISOMERIZATE
FRrtCTIONATOK
<2
flf.CUMULATOR
N-PARAFTIN
TRACT ION
Figure F4-23. Process flow diagram - liquid phase isomcrization.
Source: Reference 18
-------
--TABLE F4-59.
OPERATING CONDITIONS FGRS'ARAFFIN
ISOMERIZATION PROCESSES" -=-
Reactor Temperature, °F
Reactor Pressure, psig
Liquid Space Velocity, vol/hr/vol
Hydrogen to Oil Feed Ratio (Cs-Ce)
Hz Consumption (Cs-C6), SCF
Utilities
Fuel, 103 Btu
Electricity, kwh
Solid Bed
Systems
200-600
200-1000
1-4
1-4:1
85-160
10-50
1-2
Liquid Phase
Systems
150-250
300-500
2-3
NA1
6
NA
NA
:NA - Not Available
Source: References 1, 2, 4, 18, 20, 22
4.5.3.3 Process Emissions--
The isomerization process is a closed system. That
is, there are no process streams vented to the atmosphere. The
emission sources for this process are:
• Process heater flue gas, and
• Fugitive emissions.
Process Heater Flue Gas--Emission factors for various
pollutants from fired heaters are given in Table F4-60. A
typical solid-bed isomerization unit will require one process
heater to preheat the feed prior to entering the reactor. Heat
input requirements vary, ranging from 10,000-50', 000 Btu per bar-
rel of feed.
316
-------
TABLE F4-60.
TYPICAL EMISSIONS FROM ISOMERIZATION
UNIT PROCESS HEATERS ~~T ^
EPA Emission Factor
(lb/10} gal-oil fired)
(lb/106 scf-gas fired)
Total Emissions
(lb/10?.bbl of
fresh feed)
Oil Fired Heaters
Particulates
- Distillate oil
- Residual oil
Grade 4
Grade 5
Grade 6
Sulfur Dioxide0
- Distillate oil
- Residual oil
Sulfur Trioxidec
Carbon Monoxide
Hydrocarbons (as
Nitrogen Oxides
(as N02)
- Distillate oil
- Residual oil6
7
10
10(S)+3
U2(S)
157Q,5)r use emission
factor of 120 lb/103 gal
Based-on sulfur content of 2000 gr/106 scf
317
-------
'-£• Total emissions from the isomeriza'fTion unit heater are
also given in Table F4-60. These figures are given as pounds of
pollutant per thousand barrels of feed using a heat input of
50,000 Btu per barrel feed.
Fugitive Emissions — Fugitive emissions can occur from
a variety of sources including valves, pumps, cpmpressors,
flanges and other fittings, relief valves, and drains. Hydro-
carbon emission rates for each individual source are quite low.
However, total fugitive emissions are significant because of the
large number of such sources within the refinery.
Total estimated fugitive non-methane hydrocarbon emis-
sions from a typical butane isomerization unit- are given in
Table F4-61. The listed emission factors were determined as a
part of this program. Additional information on these emission
.factors, including a complete discussion on their derivation
and the confidence intervals for each source category, are con-
tained in Appendices B and C of this report.
Estimates for the number of sources within each source
category were developed from:
• Actual field source counts conducted during
the course of this program, and
_--"" • Counts contained in The Development of Petro-
leum Refinery Plot Plans, EPA publication
number 450/3-78-025.10
The methods used to develop estimates for the number
of each- source type have been previously described in Section
4.1.1.3 of this Appendix.
318
-------
TABLE F4-61.
ESTIMATED FUGITIVE NON-METHANE HYDROCARBON EMISSIONS
FROM A TYPICAL BUTANE ISOMERIZATION UNIT
• \( '•
Emissions
Source Type
Valves
Open-End
(Sample)
Valves
Purops (Pump
Seals)
Drains
Flanges &
Fittings
Relief Valves
Compressors
fComDresdor
Process
Stream Service
Classification
Gas/Vapor
Light Liquid
(V? > 0.1 pela 9 100'Tt
Heavy Liquid
(V? < 0.1 psla 9 100°P)
Hydrogen Service
Total
All
Light Liquid
(V? > 0.1 pal« 9 100'F)
Heavy Liquid
(VP < 0.1 pala 9 100'F)
Total
All
All
All
Hydrocarbon
HvdtoBen
Number of Sources in Process Unit
238
310
0
102
"650*
_
10(14)
fl ( C\"\
10(14)8
26a
2321a
6a
0(0)
2(4)
Source
Emission
Factor, Ib/hr
0.059
0.024
0.0005
0.018
0.005
0.25
0.046
0.070
0.00056
0.19
1.4
0.11
Estimated Total
Emissions,
Ib/hr Aj,
14.0
7.44
0.0
1.84
23.3
-
3.50
0.0
3.50
1.82
1.30
1.14
0.0
0.44
Estimated
-------
-£. The source counts listed in Table "£4^61 refer to those
sources located within the battery limits of the process. That
is, equipment located in tankage or transfer lines is not in-
cluded in these source counts.
The number of sources per unit depends on a variety of
factors including processing complexity, process type, and pro-
cessing flexibility. The source counts given in Table F4-61 are
not necessarily representative of all or even the majority of
isomerization units.
4.5.4 Hydrodealkylation
Aromatic hydrocarbons, required in the production of
plastics, fibers, rubber, detergents, insecticides, etc., are
produced within the refinery from certain petroleum fractions.
This source of aromatics was established in the early 1950's as
a result of:2 3
• The development and rapid growth of the
catalytic reforming process,
• The development of improved separation
processes for isolating high purity
aromatics, and
~~ • The insufficiency of other sources of aromatics
(such as coal tar by-products) to meet rising
demand.
The single ring aromatics produced in catalytic re-
forming-include benzene, toluene, and xylenes. ' Of these, the
demand .for toluene and m-xylene is low. However, considerable
320
-------
quantities of these hydrocarbons are produced- during the re-
^ — -^_
forming process. A similar situation exists for double ring
aromatics as the demand for naphthalene exceeds that for the
alkyl-naphthalenes. Hence, processes have been" developed to
convert low value aromatics into higher value products. One
such process is hydrodealkylation.
4.5.4.1 Process Description--
The hydrodealkylation process serves to remove alkyl
groups from aromatic rings at elevated temperatures in the
presence of hydrogen. Toluene and xylene are converted to
benzene while alkyl-naphthalenes are converted to naphthalene.
A number of commercial hydrodealkylation processes
have been developed. The reaction can be conducted either by
using a catalyst at pressures of 280-850 psig and temperatures
of 1020-1170°F, or thermally, in the absence of a catalyst at
1200-1440°F in the same pressure range.23
At present, 11 hydrodealkylation units are in opera-
tion with a total installed capacity of 29,340 bbl/day.3 Most
of these units appear to be producing benzene from feed obtained
from BTX operations.
4.5.4.2 Process Technology--
A simplified flow diagram for the combined hydrode-
alkylation of single and double ring aromatics is given in
Figure F4-24. Fresh feed and unconverted recycle materials
are mixed with hydrogen, preheated, and charged to the reactor.
The hydrogen, often from a reformer, may range in purity from
25-95" "percent. A hydrogen purification unit can be installed
to reduce hydrogen consumption.
321
-------
tN)
HAKE-UP
HYDROGEN
I
1 PURIFICATION i
I OR GENERATION .
y\
REACTORS
T
FLASH DRUM
t
T-.
UNCONVERTED PRECURSORS
L
r
I
FUU GAS
H'
r
NAPMTIIALFNE
*•* BCN7ENE
Figure F4-24. Process flow diagram - hydrodealkylation.
Source: Reference 23
-------
^- At reactor conditions, alkyl groups-axe removed from
•*£" — -C_
the aromatic base with the liberation of light gases. Non-
aromatic hydrocarbons may be cracked to lighter hydrocarbons.
In addition, sulfur in the feed, if any, is converted to HaS,
leaving the products essentially sulfur free.2^
The reactor products discharge into a high pressure
separator where hydrogen and most of the light hydrocarbons
(Ci-Cs) are removed. Liquid from the separator is stabilized
and fractionated into benzene, naphthalene, and an unconverted
recycle stream.
A variety of feedstocks may be used provided that the
aromatics concentration is sufficiently high. .Typical feeds
for benzene production are toluene and/or xylene. Xylene feeds,
however, are less economical than toluene as a larger portion
of the feed is converted to gaseous products. Hydrogen con-
sumption is also increased with xylene feeds.23
Feeds for the production of naphthalene include
heavy reformate, catalytic cracker cycle oils, and other heavy
aromatic feeds. To avoid excessive hydrogen consumption, the
aromatic components in these streams must be concentrated by
distillation, solvent extraction, or other separation processes.
4.5.4.3 Process Emissions--
Emission sources from hydrodealkylation processes
include:
• Process heater flue gas, and
.~~~~ • Fugitive emissions.
323
-------
:-~ Process Heaters--Emission factors "fbi£_various pollu-
tants from process heater are given in Table F4-62. A typical
hydrodealkylation process will require one process heater. It
is used to preheat the feed prior to its entry, into the reactor.
In addition, a process heater may be used as a-reboiler for the
main fractionation column, particularly when processing heavier
aromatic feeds.
Total emissions for the hydrodealkylation heater are
also given in Table F4-62. These figures are given as pounds
of pollutant per thousand barrels of feed using a heat input
of 290,000 Btu per barrel of fresh feed.
Fugitive Emissions—Fugitive emissions can occur from
a variety of sources including valves, pumps, compressors,
flanges and other fittings, relief valves, and drains. Hydro-
carbon emission rates for each individual source are quite low.
However, total fugitive emissions are significant because of the
large number of such sources within the refinery.
Total estimated fugitive non-methane hydrocarbon emis-
sions from a typical hydrodealkylation unit are given in Table
F4-63. The listed emission factors were determined as a part of
this program. Additional information on these emission factors,
including a complete discussion on their derivation and the con-
fidence intervals for each source category, are contained else-
where -rh this report.
Estimates for the number of sources within each source
category were developed from:
— • Actual field source counts conducted during
; the course of this program, and
324
-------
TABLE-F4-62.
TYPICAL EMISSIONS FROM HYDRODEALKYLATION
UNIT PROCESS HEATERS - Z.
EPA Emission Factor
(lb/103 gal-oil fired)
(lb/106 scf-gas fired)
Total Emissions
(lb/l&3 bbl of
fr_esh feed)
Oil Fired Heaters
Participates
- Distillate oil
- Residual oil
Grade 4
Grade 5
Grade 6
Sulfur Dioxide
- Distillate oil
- Residual oil
Sulfur Trioxidec
Carbon Monoxide
Hydrocarbons (as CHi,)'
Nitrogen Oxides
(as N02)
- Distillate oil
- Residual oil6
7
10
10(S)+3
142(S)
157(S)
2(S)
5
1
22
22+400 (N^
4.1
15
21
21(S)+6.2
294(S)
325(S)
4.1(5)
10
2.1
46
46+829(N)'
Gas Tired Beaters
Participates
Sulfur Oxides (as S02)f
Carbon Monoxide
Hydrocarbons (as CHi,)
Nitrogen Oxides (as N02)
5-15
0.6
17
3
120-230
1.4-4.1
0.17
4.7
0.83
33.1-63.5
Source: Reference n
Based on a heat input of 290,000 Btu/bbl of fresh feed with the following
fueljveatlng values: Oil - 140,000 Btu/gal; Gas - 1050 Btu/scf.
°S » wt 2 sulfur in the oil
Improper combustion may cause a significant increase in emissions
*Use this emission factor for residual oils with less- than 0,52 C"<.5). nitro-
gen content. For oil with higher nitrogen content O^O^SJ, use emission
factor of 120 lb/103 gal
Based-on sulfur content of 2000 gr/106 scf
325
-------
TABLE F4-63.
ESTIMATED FUGITIVE NON-METHANE HYDROCARBON EMISSIONS FROM
A TYPICAL HYDRODEALKYLATION UNIT
LO
• \ ! ''
Emissions
Source Type
Valve u
Open-End
(Sample)
Valves
Pumps (Pump
Seals)
Drains
Flanges 6
Fittings
Relief Valves
Compressors
(Compressor
'Seals) '
8Counted from
Estimated
CReference 10
Process
Stream Service 1
Classification
Gas/Vnpor
Light Liquid
(VP > 0.1 pels 9 100'F)
Heavy Liquid
(VP < 0.1 pela e lOO'F)
Hydrogen Service
Total
All
Light Liquid
(VP > 0.1 pela S 100'F)
Heavy Liquid
(VP < 0.1 p»la 3 100'F)
Total
All
All
All
Hydrocarbon
Hydrogen
Total
Flow Diagrams
Number of Sources
In Process Unit
« -
founts or Estimates Counts or Estimates Em:
From Radian Study
179
391
43
22.
690b
-
13(18)
K 2)
14(20)b
36b
246 3b
6b
0(0)
3(6)
jrce
Ission
Estimated Total ,
Emissions. \i
From PES Study c Factor, Ib/hr Ib/hr
116
352
0
100
~568b
10*
6(8)
0(0)
6(8)a
-
1880b
-
0(0)
2(4).
2(4)a
0.
0.
0.
0.
0.
0.
0.
0.
0.
0.
1.
0.
059
024
0005
018
005
25
046
070
00056
19
4
11
6.
8.
0.
1.
16.
2.
0.
2.
1.
0.
o.
23.
84 -
45 -
0 -
39 -
7 -
0.
00 -
0 -
00 -
2.
05 -
1.
0.
44 -
44 -
9'-
10
9
0
1
21
05
4
0
4
52
1
14
0
0
0
32
.6
.38
.022
.80
.8
.50
.092
.59
.38
.66
•<*i
.I/''"'
-------
JT • Counts contained in The Development of
Petroleum Refinery Plot Plans, EPA pub-
lication number 450/3-78-025.:°
The methods used to develop estimates- for each source
type have been previously discussed in Section 4.1,1.3 of this
Appendix.
The source counts listed in Table F4-63 refer to those
sources located within the battery limits of the process. That
is, equipment located in tankage or transfer lines is not in-
cluded in these source counts.
The number of sources per unit depends on a variety of
factors including processing complexity, process types, and
processing flexibility. Thus, the source counts given in Table
F4-62 may not be representative of all or even the majority of
hydrodealkylation units.
4.5.5 References for Section 4.5
Knight, W. N. N., and M. L. Peniston-Bird. Modern Petro-
leum Technology. Applied Science Publishers, Ltd., Essex,
England, 1975. 293 pp.
Gary, James H., and Glenn E. Handwerk. Petroleum Refining,
Technology and Economics. Dekker, New York, 1975.
Cantrell, Ailleen. Annual Refining Survey. Oil and Gas J.,
March 26, 1979.
Hydrocarbon Processing Refining Process Handbook, 57(9),
September 1978.
Hoffman. H. L. A Bigger Role for Cat Reforming. Hydro.
Troc. 50(2) 85-88, February 1971. ^_
Lo*vell, L. M. , T. M. Moore, and R. D. Petersen. Catalytic
Reforming. Am. Chem. Soc. Div. Pet. Chem. Prepr. 17(3),
July 1972.
327
-------
7. Wljat's New in Cat Reforming? The PetroTeife Publishing
Company, Tulsa, Oklahoma, Undated.
8. Pollitzer, E. L. Continuous Platforming. Plat. Met. Rev.,
20(1): 2-6, 1976.
9. Cha, Bernard J., Roland Huin, Hugo Van Landeghem, and Andre
Vidal. Regenerate Reformers Continuously. Hydro. Proc.,
52(5): 98-101, May, 1973.
10. U.S. Environmental Protection Agency, Office of Air Quality
Planning and Standards. Development of Petroleum Refinery
Plot Plans. EPA-450/3-78-025. Research Triangle Park,
North Carolina, 1978.
11. Environmental Protection Agency. Compilation of Air Pol-
lutant Emission Factors. Third Edition, Supplement No. 8.
AP-42. Research Triangle Park, N.C., May 1978.
12. Dickerman, J. C. , et al. Industrial Process Profiles for
Environmental Use: Chapter 3, Petroleum Refining Industry.
Final Report. EPA-600/2-77-023C, Radian Corporation,
Austin, Texas, January 1977.
13. Air Pollution Control District, County of Los Angeles;
State of California, Department of Public Health; U.S.
Department of Health, Education and Welfare; Western Oil
and Gas Association. Emissions to the Atmosphere from
Eight Miscellaneous Sources in Oil Refineries. June 1958.
14. Lafferty, W. L., Jr., and R. W. Stokeld. Alkylation and
Isomerization. Adv. Chem. Ser., 103: 130-49, 1971.
15. McGovern, L. J. Developments in Commercial Alkylation of
Isobutane the Past 25 Years. ACS Div. Pet. Chem. Prep.,
17(3): B19-B34, 1972.
16. Anderson, R. F. Changes Keep HF Alkylation Up-to-Date.
Oil and Gas J., 72(6): 78-82, February 11-, 1974.
17. Fenske. E. R. The Modern Hydrogen Fluoride Alkylation
Unit. 'CEP Symposium Series, 57(34), AIChE, N.Y., Undated.
18. Ewing, Robert C. Making Isoparaff ins. Oi-1 and Gas J. ,
69(33): 61-72, August 16, 1971.
19. Asselin, G. F. Isomerization of Paraffins-- Am. Chem. Soc.
-Div. Pet. Chem. Prepr., 17(3), July 1972; August 27,
1972; September 1, 1972.
328
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20. ClJ&pey, Nicholas P. Isomerization Is In2.z^ Lead Additives
Go Out. Chem. Eng., 78(12): 24-46, May 31, 1971.
21. Bour, Georges, C. P. Schwoerer, and G. F. Asseiin. Penex
Unit Peps Up SR Gasoline. Oil and Gas J. , 68(43): 57-61,
October 26, 1970.
22. Hydrodealkylation Processes. Ind. and Eng. Chem.,. 54(12):
28-33, 1962.
23. Sangal, M. L. , K. M. Murad, R. K. Niyogi, and K. K.
Bhattacharyya. Production of Aromatics from Petroleum
Sources. J. of Sci. Ind. Res., 31(5): 260-4, 1972.
329
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4.6 ~; GAS PROCESSING - ^
Gas processing recovers various hydrocarbons as pure
products or as mixtures of specified compositions. The prod-
ucts of gas processing may be fuel gas, methane, ethane, pro-
pane, propylene, normal and isobutane, butylenes, normal and
isopentane, amylene, and/or a light naphtha. These products
may be used in other refinery processes, used as gasoline blend-
ing components, or sold.
The feed to gas processing units is from crude dis-
tillation, catalytic reforming, catalytic cracking, hydrocrack-
ing, thermal cracking, and to a lesser extent, hydrodesulfuriza-
tion. Many refineries "spike" the crude with natural gas
liquids to increase the yield of gaseous products. Major units
include acid gas removal, dehydration, and separation.
4.6.1 Gas Treating/Cleaning
4.6.1.1 Acid Gas Removal--
The acid gas removal unit removes hydrogen sulfide
from hydrocarbon gases, usually by absorption in an aqueous,
regenerative sorbent. C02 and/or mercaptans may also be re-
moved, depending on the process used.
. ^r A number of acid gas removal processes are available,
distinguished primarily by the regenerative sorbent used.
Amine-based sorbents are most commonly used.
The feed to a typical unit is contacted with the sor-
bent, _such as diethanolamine, in an absorption--column to selec-
tively absorb H2S from the hydrocarbon gases. Hydrogen sulfide
330
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is then,-removed from the sorbent in a regeneration step. The
products are a sweet hydrocarbon gas and a concentrated hydrogen
sulfide stream. The hydrogen sulfide stream is normally routed
to a sulfur plant for recovery of its sulfur co~htent. Alterna-
tively, the sulfide gas may be flared to produce the less toxic
sulfur oxides.
Acidic gases can also be absorbed into an aqueous
caustic soda solution. In this case, the spent absorbent is
periodically dumped.
Process Conditions—A typical absorber operates at a
pressure of about 150 psi and a temperature of about 100°F.
Pressure and temperature may, in some instances, be significantly
higher.
Emissions--If a regenerative sorbent"system is used in
conjunction with a sulfur recovery unit, only fugitive emissions
are produced. If the hydrogen sulfide stream is flared, sulfur
oxide emissions are produced. If caustic soda is used, a liquid
waste stream of the spent absorbent is produced.
4.6.1.2 Sulfur Recovery--(See Section 4.7.5)
4.6.1.3 Dehydration--
. -- Dehydration removes water from the gas after the acid
gas removal step. The required content is specified as the dew
point, the temperature at which the water begins to condense.
Excess water may be removed by refrigeration, absorp-
tion, _p_r adsorption. Refrigeration processes decrease the tem-
to below the required dew point; condensed moisture is collected
for disposal.
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.—~ Absorption processes allow the mois-t gas to flow over
a hygroscopic material such as di- or triethylene glycol. Two
to five kilograms of glycol are circulated per kilogram of water
to be removed. ~
Solid dissicants such as silica gel or alumina are
used in adsorption processes. When the last traces of water,
higher hydrocarbons, dissolved elemental sulfur, and residual
sulfur compounds must be removed, a molecular sieve bed may be
used. Beds are regenerated with hot gas.
Process Conditions-- Temperature and pressure are in-
terdependent in condensation processes. For example, if the
required dew point is 50°F at 135 psig and the.best available
cooling is 80°F, the pressure will be 460 psig.
For absorption processes using di- or triethylene
glycol, absorption temperatures must be kept below the glycol"s
decomposition temperature (327°F for DEC, 405°F for TEG). Dew
point depressions of 40° to 50°F can be obtained by using DEC
at atmospheric pressure and reboiler temperatures of 290° to
320°F. Temperatures in the regenerator, where water is sepa-
rated from the glycol, usually range from 375° to 400°F. The
pH is controlled at 6.0. to 7.5: a low pH accelerates decomposi-
tion of the glycol.
==- Regeneration temperatures for solid dessicants are
480° to 500°F.
Utilities—A glycol absorption process requires about
0.1 percent of the fuel produced.
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~ Emissions--An estimated 0.1 gallon "ojf^triethylene
glycol per 10 ft3 of gas processed is emitted by a glycol
absorption unit in vented water vapor. Water contaminated with
glycol may be vented as steam or it may be disposed of as a
liquid.
4.6.2 Product Separation/LPG Production
Refinery gas is sometimes used without separation into
its dry gas and liquid components, especially if it is to be
used as fuel gas. However, the gas is often separated into its
components in a gas separation plant. This separation is usually
accomplished by contacting the gas with an absorber oil. Re-
frigerated absorption, refrigeration, or adsorp-tion may be used
when a separate methane stream is desired.
In the oil absorption process, the gas is contacted
with an absorber oil in a packed or bubble tray column. Propane
and heavier hydrocarbons are absorbed by the oil while most of
the methane and ethane pass through the absorber. The enriched
absorber oil is then taken to a stripper where the absorbed
propane and heavier compounds are stripped from the oil. If the
feed is refrigerated to a -40"F dew point before entering the
unit, all hydrocarbons except methane are absorbed.
In the refrigeration process, the gas is first dried
with molecular sieve beds to a dew point of -150°F or less. It
is then cooled in a heat exchanger to -25°F. Condensed hydro-
carbons are removed in a gas-liquid separator. The gas from
this separator is passed through a second separator at -135°F.
Liquids^ from the separators are fed to a series of distillation
columns^ where methane, ethane, propane, butane^'and other prod-
ucts are recovered.
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^- An activated carbon bed adsorbs alt hydrocarbons ex-
cept methane. The bed is regenerated with heat and steam; the
resulting hydrocarbon vapor is condensed and the water separated.
The resulting hydrocarbon product is then fractionated into its
various components.
Process Conditions—Pressure in an oil absorber may
be as high as 400 psi, but is usually lower. Inlet gas and oil
temperatures are 90° to 100°F.
Emissions--Fugitive emissions from leaking pumps,
valves, compressors, and other fittings are the only emissions
from product separation.
4.6.3 References for Section 4.6
1. Campbell, John M. Absorption and Fractionation Funda-
mentals. Gas Conditioning and Processing, John M.
Campbell, Norman, Oklahoma, 1970.
2. Cavanaugh, E.G., et al. Atmospheric Pollution Potential
from Fossil Fuel Resource Extraction, On-Site Processing,
and Transportation. Final Report. EPA-600/2-76-064,
Radian Corporation, Austin, Texas, March 1976.
3. Cotterlaz-Rennaz. New French Gas Cooler Recovers 120 BPD
Gasoline. World Oil, 177(2): 57-59, 1973.
4. Eckerson, B. A. , and A. L. Johnson. Natural Gas and Natural
Gas Liquids. In.- Surface Operations in Petroleum Production.
George V. Chilingar and Carrol M. Beeson, -eds. American
Els'evier Publishing Company, Inc., New York, 1969.
5. Ecology Audits, Inc. Sulfur Compound Emissions of the
Petroleum Production Industry. EPA-650/2-75-030, Dallas,
Texas, 1974.
6. Natural Gas Producers Suppliers Association. Gas Dehydration,
from Dehydration and Treating. Engineering, Data Book. 1972.
334
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7. HSmmerschmidt, E. G., K. R. Knapp, and G-. ^L,. Perskin.
Gas Hydrates and Gas Dehydration. Gas Engineers Handbook.
Industrial Press, New York, 1969.
8. Houghton, J. and J. D. McLay. Turboexpanders Aid Con-
densate Recovery. Oil and Gas J., 71(10): 76-79, March 5,
1973.
9. Hydrocarbon Processing Refining Processes Handbook. Hydro-
carbon Proc., 53(9), September 1974.
10. Nack, H., et al. Development of an Approach to Identifica-
tion of Emerging Technology and Demonstration Opportunities.
EPA-650/2-74-048, Battelle-Columbus Labs., Columbus, Ohio,
1974.
11. NG/LNG/SNG Handbook. Hydrocarbon Processing., 52(4), April
1973.
12. Patterson, E. 0., Jr. Get Low Dewpoints with Solid
Dessicants. Oil and Gas Journal 67(9): 108-109, 1969.
13. Petroleum Extension Service. Field Handling of Natural Gas,
3rd ed. The University of Texas at Austin, Austin, Texas,
1972.
14. Processes Research, Inc. Screening Report, Crude Oil and
Natural Gas Production Processes, PB-222718, Cincinnati,
Ohio, 1972.
15. Radian Corporation. A Program to Investigate Various
Factors in Refinery Siting. Final Report. Contract No.
EQC 319, Austin, Texas, 1974. (Unpublished)
16. Hobson, G. D., ed. Modern Petroleum Technology. Essex,
England. Applied Science Publishers, Ltd., on behalf of
The Institute of Petroleum, Great Britain, 1973.
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4.7 OTHER PROCESSES
•<" — i_
4.7.1 Asphalt Processing/Production
Asphalt is produced as the bottoms from vacuum distilla-
tion, as discussed in Section 4.1.2. Any lube oil present can be
removed by deasphalting, as described in Section 4.7.2.
4.7.1.1 Process Description--
Asphalt blowing oxidizes the asphalt to increase its
melting temperature and hardness. Both batch and continuous
processes are employed. Fresh feed and recycle are heated to
approximately 500°F and charged to a vertical vessel. Pressur-
ized, preheated air (390-590°F) is charged into the bottom of
the vessel through a sparger. The reaction is exothermic and
quench steam is sometimes required for temperature control.
Excess air is minimized to avoid combustion of exhaust fumes.
In some cases, ferric chloride or phosphorus pentoxide are used
as catalysts to increase the reaction rate and impart special
characteristics to the asphalt. Utilities required for asphalt
blowing include 5,000 to 10,000 Btu/bbl heat required for heating
the asphalt to the reaction temperature and one kWh/bbl for air
compression.
4.7.1.2 Atmospheric Emissions--
_.__ The quantity of hydrocarbon emissions.from asphalt-
blowing units should be relatively small since the asphalt is
distilled at high temperatures before reaching the air-blowing
process. Available data indicate that uncontrolled emissions
amount to 60 pounds per ton of asphalt.1 The operating condi-
tions are favorable for the production of extreffiely undesirable
polynuc-lear aromatics.
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^ In some refineries, air-blown brigh£q£iing units have
*r - . ~-
been replaced by vessels packed with solid absorbents. These
vessels have no hydrocarbon emissions other than fugitive emis-
sions. ~~
4.7.2 Lube Oil Processing/Production
Lube oil stock is produced as the 700° to 1000°F frac-
tion of the residium from vacuum distillation. For better con-
trol of lube oil processing, the vacuum distillation unit is
usually built as part of the lube oil plant.2 Operation of the
vacuum distillation unit is discussed in Section 4.1.2.
Procedures for processing the lube oil stock into
specific products vary greatly with the individual refinery. It
is not possible to present any one flow scheme as representative
of the industry. Processing procedures can, however, be cate-
gorized into four groups: deasphalting, treating, dewaxing, and
finishing.
Each of these processes is closed to the atmosphere.
Except for hydrotreating, there are no emissions other than fugi-
tive emissions from valves, flanges, etc., and emissions from
process heaters. Because of the low volatility of the charge
stock, fugitive emissions of the material are expected to be
generally low. However, more volatile materials may be used in
certaia-treating operations and fugitive emissions of these
materials can be significant. With hydrotreating, there are
emissions, particularly CO, associated with the periodic re-
generation of the catalyst.
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4. 7 . 2. £-' Deasphalting- - '~~^._
A very heavy oil (brightstock) can be_ produced from
vacuum residues by extraction with propane. The process, called
propane deasphalting, is particularly effective because propane
is a very selective solvent. At temperatures from. 104°F to
140°F, paraffins are quite soluble in propane, but high molec-
ular weight asphaltic and resinous compounds precipitate. The
extract phase contains about 15-20 weight percent oil dissolved
in propane. The raffinate is a colloid or emulsion of precipi-
tated asphalt material in propane.
The major equipment for the above process includes a
heat exchanging system, an extractor, and a propane recovery
system. Propane recovery is usually in two stages. The first
stage is operated at an elevated pressure; and recovered propane
is condensed by air or water cooling. In the second stage, the
propane is stripped at atmospheric pressure. The stripping steam
is then condensed and the propane vapor compressed and condensed.
A relatively large amount of propane is required for the produc-
tion of commercial quantities of lube oil.
In special cases, propane can be used to separate a
lighter oil fraction (SAE50), brightstock, and hard asphalt.
The fractionating property of propane is used for this separa-
tion. For this fractionation, a propane deasphalting unit con-
sists of. two units in series. The raffinate from the first unit,
along with more solvent, is sent to the second unit which oper-
ates at a lower temperature. The second unit splits the raffin-
ate into a heavy oil and hard asphalt.
Old propane deasphalting units used a~~"mixer-settler
system for contacting the vacuum residue with the propane, but,
338
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because;r.bf the high viscosity of the residue,-t&e contact was
poor. A counter-current tower with perforated baffles is now
widely used for this purpose. A rotating disc contactor has
also been used successfully, but it has a relatively complicated
mechanical seal on the rotating shaft.
4.7.2.2 Treating--
There are several treating processes used to improve
the characteristics of the lube oil. Of particular interest are
the viscosity index, the color, and the carbon residue content.
The two most popular treating methods are phenol extraction and
furfural treating; the choice between these two is often one of
solvent availability. Hydro treating has also been used.
Phenol Extraction--Aromatic and naphthenic hydrocarbons
are particularly soluble in phenol. Because of this, a relatively
low solvent to oil ratio is required. In fact, some hydrocarbons
dissolve in phenol so readily that water must be injected to
control this solubility. The degree of extraction and the
selectivity of the process are controlled by several variables:
solvent to oil ratio, extraction temperature, temperature gradient,
and water injection rate.
In the process, oil is extracted with phenol in a
counter-current extractor; water is injected into the extract
phase as it leaves the extraction zone. Hydrocarbons which
separate from the extract upon addition of water are returned to
the extraction zone. This intensive mass transfer operation
results in a sharp separation of the extract products from the
raffinate.
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±- Hydrocarbons are less soluble in pbeflol at lower tem-
peratures; therefore reducing the temperature of the extract
phase as it leaves the extraction zone has a similar effect to
that of adding water. In this case, the extraction tower is
operated with a temperature gradient. Phenol entering the
counter-current extractor is at a higher temperature than the
oil, therefore the raffinate phase leaving the extraction has a
higher temperature than the extract.
Water is partly miscible with phenol and also forms an
azeotrope with the phenol, therefore it must be removed in a
drier as a vapor phase phenol-water azeotrope. The phenol is then
absorbed in the oil entering the process.
A phenol extraction unit consists of an extraction stage
and a phenol recovery section. The phenol recovery section
includes flash towers and vacuum steam strippers.
The counter-current packed tower or a tray tower has
been used for phenol extraction since the process was first used
for the treatment of lube oil. Modern units use either rotating
disc contactors or Podbielniak extractors.
The Podbielniak extractor is a horizontal centrifuge in
which counter-current flow is accomplished by centrifugal force.
The extractor is highly efficient and has a much smaller volume
than do.es an extraction tower. The small volume makes possible
quick changes in feedstock necessitated by market demands. This
extractor is, however, a very complicated piece of equipment.
Scale and dirt in the feedstock can cause vibration if unevenly
distributed across the cross-section of the rotating drum.
•-e£
—;" The rotating disc contactor (RDC) consists of a series
of compartments formed by stator rings and rotating discs. The
340
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heavy ghase falls onto the rotating disc and" 3.£[_ thrown tangen-
tially against the incoming light phase. An RDC may be one
third the height of a packed tower with equal extractive ability.
The smaller volume of the RDC makes possible quick change-overs
of feedstock.
Furfural Treating--Furfural treating is similar to
phenol extraction. Two main differences are that (1) a larger
solvent to oil ratio is required and (2) there is no need for
water injection to reduce the solubility of certain hydrocarbons.
Furfural can be easily oxidized to organic acids which
are extremely corrosive, therefore some furfural treating units
have been designed with charge stock deaerators to keep air out
of the system. Extraction was originally accomplished in packed
towers, but since the 1950"s the rotating disc contactor has
become popular.
Hydrotreating-- Hydrotreating is used for viscosity
index improvement, desulfurization, denitrogenation, demetalliza-
tion, removal of gum forming compounds, and color improvement.
The oil feed is mixed with make-up and recycle hydrogen
and charged to a fixed catalyst bed reactor. Reactor effluent
flows through high and low pressure separators for removal of
hydrogen for recycle and light ends, respectively. The product
is then^steam stripped to remove any remaining impurities.
The feed to a hydrotreating unit can be either solvent
refined lube oils and waxes or raw distillates and deasphalted
oils. . _Hydrogen requirements vary from 100 to 200 cubic feet H2
per barrel of oil. The catalyst generally has a cobalt or nickel-
molybdenum base.
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^- The reactor operates at 600 to 750°E ^nd 500 to 700
psi. Utilities for the treatment of one barrel~~bf oil are 2-5
kwh electricity, 15 to 30 Ib steam, and 35,000 to 140,000 Btu
heater fuel. ~~
Other Treating Processes—Treatment of lube oil with
sulfuric acid is a simple matter of contacting the oil with the
acid and then separating the resulting acid sludge from the re-
fined oil by clay filtration. This process was once used more
extensively, but is seldom used now because of operating diffi-
culties and sludge disposal problems.
Duo-sol, a process using propane and cresilic acid for
simultaneous deasphalting and treating, was once used for about
20 percent of the lube oil produced. It is seldom used now be-
cause of increased quality requirements for lube oil products.
Other solvents which have been suggested for lube oil
treating are liquid S02, a liquid S02/benzene mixture, a mixture
of cresols, a mixture of phenol and cresols, nitrobenzene, ani-
line, and dichloroethylether. Some of these have been used in
commercial plants.
4.7.2.3 Dewaxing--
Dewaxing is the most difficult part of lube oil manu-
facture.,- The dewaxing process removes wax from- lube oils to
improve the low temperature fluidity characteristics of the oil.
The oil is contacted with solvent and chilled, causing the wax
to precipitate. The precipitated wax is separated from the mix-
ture by filtration or centrifuging. The dewaxed oil and solvent
are separated by distillation and steam stripping. Solvent is
recycled. The wax, usually containing at least"10 percent oil,
is solv'ent treated again under different conditions to obtain a
342
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deoileji-wax product of the desired specifications. Refrigera-
" *?' —- -^.
tion and filtration are used to recover the wax and solvent.
The most widely used solvent for oilUewaxing and wax
deoiling is methyl ethyl ketone (MEK) or a mixture of MEK and
toluene or benzene. Both operations are frequently combined in
one unit using a MEK solvent. Other solvents used in oil dewax-
ing and wax deoiling are methyl butyl ketone, either alone or
mixed with toluene or benzene, and propane, acetone, and chlori-
nated hydrocarbons.
Oil dewaxing and wax deoiling processes are major
energy consumers because of refrigeration and filtration re-
quirements. Two to ten kwh of electricity and.100-400 pounds
of steam are required per barrel of oil.
4.7.2.4 Finishing —
The color of a product can be quickly deteriorated by
the presence of traces of resinous materials and chemically
active compounds. There are several finishing methods for re-
moving these compounds.
The compounds can be absorbed by contacting the oil
with various types of mineral clays, activated earth, or arti-
ficial absorbents. The oil is either mixed with the absorbent
or percolated through a long column packed with the absorbent.
Absorption occurs at high temperatures, about 600°F. The ab-
sorbent is then removed from the oil by filtration. Clay ab-
sorption is no longer used in large-scale lube -oil manufacture
because of problems with disposal of the spent absorbent.
• *z- -
"~7 Hydrotreating is also used for finishing lube oils
(hydro'finishing) . Hydrofinishing effectively removes nitrogen
343
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compounds which cause the oil to darken and to be unstable. Re-
' ^ • .^ -
movel of nitrogen compounds usually requires considerably higher
pressures than do middle distillate desulfurizers, therefore the
sulfur content of the oil is also reduced. Hydrofinishing also
removes oxygen very efficiently.
In hydrofinishing, the charge stock is combined with
hydrogen gas, heated in a furnace, and passed through a hydro-
finishing reactor. The reactor effluent is cooled and the liquid
and vapor phases are separated. The vapor phase is recycled.
The liquid is sent to a stripper where H2S, NH3 and light hydro-
carbons are removed.
4.7.3 Blending Operations
Refinery blending operations involve the mixing of
various components to achieve a product of desired characteris-
tics. The most common blending operation in petroleum refining
is the final step in gasoline manufacturing. Gasoline components
such as catalytic gasoline, reformate, alkylate, isomerate, bu-
tane, lead, and dye are mixed in proportions required to meet
gasoline-marketing specifications.
4.7.3.1 Process Description and Technology--
There are two methods of blending: batch and in-line.
Batch hiending is accomplished in a blending tank (or tanks) in-
to which each component is added individually. Mixing is con-
tinued until a homogeneous mixture of the desired properties is
produced. The final blend is routed to storage-tanks to await
transfer out of the refinery or pumped directly to transportation
facilities.
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'• £. Agitation in the blending tank is accomplished either
by an external circulation loop (or loops) or by internal im-
pellers powered by external motors. The impeller shafts are
sealed in the same ways as are rotating pump s.hafts. A special
case is the blending of butane into gasoline, wherein liquid
butane is sometimes charged through a sparger ring.in the bottom
of the blending tank.
In-line blending can be either partial or continuous.
Partial in-line blending involves simultaneous combination of
stock components in a mixing manifold. Final additions and ad-
justments are made downstream or in a storage tank.
Continuous in-line blending involves -continuous and
simultaneous blending of all stock components and additives in
a mixing manifold. Each component stream is controlled auto-
matically by a feedback control loop; the entire control system
is often under computer guidance. There is no blending tank,
and storage capacity is often minimized by direct discharge of
blended products to transportation facilities or pipeline.
4.7.3.2 Atmospheric Emissions--
Agitation in batch blending operations increases the
evaporation of lighter components. Thus, fugitive losses from
batch blending tanks are generally greater than those from simi-
lar qu.fescent storage tanks. Emissions from in-line blending
are limited to fugitive leaks from valves, flanges, and other
process equipment.
.._ Control technology for batch blending^ operations in-
cludesHfloating roofs on blending tanks and replacement of batch
operations by in-line blenders. Hydrocarbon leaks from in-line
345
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blendip-g systems can be reduced by proper inspection and main-
tenance of valve stem seals, flange gaskets, and pump seals.
4.7.4 Hydrogen Production
Hydrogen is a by-product of several refining processes.
However, a refinery with a large distillate hydrotreater or gas
oil hydrocracker requires additional high purity hydrogen. It
is estimated that by 1980, slightly less than 40 percent of the
hydrogen used in refineries will be manufactured.3
A steam-hydrocarbon reforming process is commonly used
for hydrogen production. However, the light hydrocarbons used
as fuel for this process are more economically- used in other
processes such as alkylation and catalytic reforming. Therefore,
steam-hydrocarbon reforming will probably be replaced by partial
oxidation of heavy oils. The choice between the two processes
will depend on the cost and availability of raw materials.u
Both processes are described here.
4.7.4.1 Steam-Hydrocarbon Reforming--
Process Technology—Either a light natural gas-type
product or a heavier hydrocarbon mixture such as naphtha can be
used as feedstock for the steam-hydrocarbon reforming process.
The feedstock is first desulfurized to prevent catalyst deactiva-
tion. -^Desulfurization may be accomplished by adsorption on
activated carbon at ambient temperatures, high temperature
reaction with zinc oxide, or catalytic hydrogeneration followed
by reaction with zinc oxide.
-^ After desulfurization, the feedstock"-and superheated
steam are catalytically reacted in a high temperature reactor
346
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(reformer) at about 1400 to 1600°F. A nickeT; fetalyst is used
for the light hydrocarbon feedstock; a special catalyst is used
for the heavier hydrocarbons. 5 __
The reformed gas contains hydrogen, -carbon monoxide,
carbon dioxide and excess steam. After the gas is. cooled, the
carbon monoxide is reacted with steam in a shift converter to
produce carbon dioxide and more hydrogen. An iron catalyst is
used.
The carbon dioxide and hydrogen-rich gas is then cooled
and scrubbed to remove practically all of the carbon dioxide.
Carbon dioxide may be removed by one of several methods, depending
on the cost of utilities. Two such methods are monoethanolamine
(MEA) scrubbing and the more thermally efficient Benfield acti-
vated hot carbonate process. The remaining carbon dioxide and
carbon monoxide are removed by heating the gas and passing it
through a nickel base methanation catalyst. Here the carbon
dioxide and carbon monoxide react with hydrogen to form methane.
Typically, about 280,000 SCF of natural gas or 60 barrels of
naphtha are required for the production of one million SCF of
^~
hydrogen. A flow diagram of the steam-hydrocarbon reforming
process is given in Figure F4-25.
Major equipment for the process includes the reformer
and the shift converter. The reformer consists of catalyst-filled
tubes..—The converter has two catalyst beds, the first containing
a high temperature catalyst and the second a low temperature
catalyst. The gas is cooled between beds.
Emissions--Emissions from steam-hydrocarbon reforming
hydrogen units include:
347
-------
FEED
00
SHIFT
CONVERTER
HEAT
RECOVERY
DESUL-
FURIZER
. , __ yf
C02
t 1
H?
C02
ABSORBER
SOLUTION
REGENERATOR
HYDROGEN
METHANATOR
Figure F4-25,
Hydrogen production by steam
reforming.
-------
' • ~f. • Emissions of process heater f itieF gas , and
• Fugitive emissions.
Emissions of various pollutants from process heaters are dis-
cussed in Section 4.1.
Fugitive Emissions — Fugitive emissions can occur from
a variety of sources including valves, pumps, compressors, flanges
and other fittings, relief valves, and drains. Hydrocarbon emis-
sion rates for each individual source are quite low. However,
total fugitive emissions are significant because of the large
number of such sources within the refinery.
Total estimated fugitive non-methane hydrocarbon emis-
sions from a typical hydrogen production unit are given in Table
F4-64. The listed emission factors were determined as a part of
this program. Additional information on these emission factors,
including a complete discussion on their derivation and the con-
fidence intervals for each source category, are contained in
Appendices B and C of this report.
Estimates for the number of sources within each source
category were developed from:
• Actual field source counts conducted during
^~ the course of this program, and"
• Counts contained in The Development of Petro-
leum Refinery Plot Plans, EPA publication
number 450/3-78-025.7
: The methods used to develop estimates for each source
type are discussed in Section 4.1.1.3 of this Appendix.
349
-------
TABLE F4-64.
ESTIMATED FUGITIVE NON-METHANE HYDROCARBON EMISSIONS FROM
A TYPICAL HYDROGEN PRODUCTION UNIT
. . u . /• •
Emissions
Source Type
Valves
Open-End
(Sample)
Valves
Purapa (Punp
Seals)
Draina
Flanges &
Fittings
Relief Valves
Compressors
(Compressor
Seals)
Pstycesa
Stream Service
Classification
Gas/Vapor
Light Liquid
(VP > 0.1 p.iia 9 100'r)
Heavy Liquid
(vr < o.i pgta 9 loo'r)
Hydrogen Service
Total
All
Light Liquid
(VP > 0.1 p«l« 9 100'P)
Heavy Liquid
(V? < 0.1 p»U 8 100'P)
Total
All
All
All
Hydrocarbon
Hydrogen
Total
Number of Sources in Process Unit
Counts or Estimates
From Radian Study
19
93
62
8
182™
—
3(4)
2(3)
5(7)a
17C
635a
6C
0(0)
3(6)
3(6)c
Counts or Estimates
From PES Study
244
2101
0
122
2467C
46b
2(3)
0(0)
2(3)b
—
8225C
~
0(0)
3(6')b
Source
Emission
Factor, Ib/hr
0.059
0.024
0.0005
0.018
0.005
0.25
0.46
0.070
0.00056
0.19
0.4
0.11 ,
Estimated Total
Emissions, .'
ib/hr •';•
1.12 - 14.4
2.23 - 50.4
0.0 - 0.031
0.14 - 2.20
3.49 - 67.0
0.23
0.75 - 1.0
0.0 - 0.138
0.75 - 1.14
1.19
0.36 - 4.61
1.14
0.0 ,
0.66; >
/ Oy6ft'fl- ,
7.82 - 7.5.6
Physically Counted
Counted From Flow Diagrams
Estimated
Reference IB
-------
~" The source counts given in Table F4--6& refer to those
sources located within the battery limits of the process. That
is, equipment located in tankage or transfer lines is not in-
cluded in these source counts.
The number of sources per unit depends on a variety
of factors including processing complexity, process type, and
processing flexibility. Hence, the number of sources shown in
Table F4-64 are not necessarily representative of all or even
the majority of hydrogen production units.
The estimated composition of fugitive non-methane
hydrocarbon emissions from a steam reformer hydrogen unit utiliz-
ing naphtha as the feedstock is given in Table-F4-65. These
estimates indicate the types of hydrocarbons contributed by
sources on a number of streams within the process unit. Addi-
tional information on these estimates may be found in Section
2.4 of Appendix D.
4.7.4.2 Partial Oxidation (Texaco Synthesis Gas Generation
Process)--
Process Technology-- Large quantities of hydrogen can
be produced economically from high sulfur residual petroleum
stocks by the partial oxidation process. The hydrogen can be
generated at any selected pressure up to 160 atmospheres. The
use of .^fuel-fired heaters is minimized; the hydrogen sulfide
produced can be recovered in the sulfur recovery unit.5
The residual fuel oil is first partially combusted with
oxygen in the presence of steam. The combustion is exothermic
and non^-catalytic. The synthesis gas produced-by the combustion
is composed primarily of hydrogen and carbon monoxide, with small
351
-------
TABLE F4-65.
ESTIMATED COMPOSITION OF FUGITIVE NON-METHANE HYDROCARBON EMISSIONS
FROM A HYDROGEN PRODUCTION UNIT UTILIZING NAPHTHA AS A FEEDSTOCK
Streams
Fuel
Gas
Straight Run
Naphtha
LPG
H? Recycle
Gas
Totals
Ui
ro
Estimated percentage of
emissions attributed to each
stream - wt %
Weighted contribution of each
component to unit emissions
p praw
Benzene
Toluene
Ethylbenzcnc
Xylenes
Other Alkylbenzenes
Naphthalene
Anthracene
Biphenyl
Other Polynuclear Aromatics
. ,1 ' «
n-llexane
Other Alkanes
Olefin
Cycloalhanes
Hydrogen
38
19
19
24
100%
0
0
0
0
0
0
0
0
s 0
0
349600
22800
0
7600
48
498
169
308
3150
278
1
119
2847
i
7379
94926
0
80277
0
0
0
0
0
0
0
0
0
0
i
0
190000
0
0
0
0
0
0
0
0
0
0
0
0
0
156000
0
0
84000
48
490
169
308
3150
278
1
119
2847
t
7379
790526
22800
80277
91600
1,000,000
-------
amount^'of carbon dioxide and minor quantiti'es ?bf hydrogen sul-
' fide, methane, nitrogen, argon, carbonyl sulfide, and soot.
The majority of the soot in the synthesis gas is re-
moved in a water quench tower. A selected soo-t level is main-
tained in the circulating water system by withdrawing a portion
of the water and contacting it with naphtha. The carbon is
transferred to the naphtha phase; the water and naphtha are
separated in a decanter and the water returned to the circulat-
ing system. The naphtha-carbon mixture is then mixed with a
portion of the residual oil-feed to the generator. The naphtha
is recovered by distillation and the soot is transferred to the
generator feed.
The gas product from the combustion chamber exits at
about 2000 to 2800°F. It is cooled with hot water and simul-
taneously saturated with steam. This steam-saturated gas stream
is fed to a shift converter where the carbon monoxide is cata-
lytically converted to carbon dioxide and additional hydrogen.
Gas from the shift converter is primarily hydrogen and
carbon dioxide with lesser amounts of carbon monoxide, methane,
hydrogen sulfide, argon, and nitrogen. The Rectisol process,
which uses methane at temperatures of -50 to -60°F, effectively
separates a pure carbon dioxide stream with less than 1 to 5 ppm
hydrogen sulfide and a 30 to 40 percent hydrogen sulfide stream
which :rs sent to the sulfur recovery unit. The remaining hydro-
gen is more than 95 to 98 percent pure. Removal of carbon mon-
oxide, perhaps with copper liquor, raises the purity of the
hydrogen to 97 to 99 percent.
-= A flow diagram of the partial oxidation process is
given as Figure F4-26. Fuel and utility requirements for the
353
-------
u>
WATER
Figure F4-26. Hydrogen production by partial oxidation.
-------
production of one million SCF 100 percent hydrogen are given as
Table F4-66.
TABLE F4-66. UTILITY REQUIREMENTS FOR THE'
PRODUCTION OF 1,000,000 SCF
100 PERCENT H2
Electric Power, kwh 1,700
Cooling Water, Gal 90,000
Deionized Water, Gal , 450
Boiler Feed Water, Gal 10,350
Low Sulfur Fuel, Bbl 0.8
Emissions--No specific information was available on
emissions from the partial oxidation method. It is assumed
that the only emissions are those from process heaters and the
fugitive emissions from valves, flanges, etc. Emissions from
process heaters are discussed in Section 4.1.
4.7.5 Sulfur Recovery
A sulfur recovery plant convents hydrogen sulfide to
elemental sulfur. The Claus process is the accepted method for
sulfur recovery in the refining industry; it is assumed to be
used by all major refiners.
C^ Process Technology—Process types and process flow
diagrams for the Claus process are given in Section 3.2.1.3 of
Appendix E. Raw materials for the unit include H2S from the
acid g-as removal plant and H2S from sour water stripper systems.
The amount of sulfur reaching the sulfur recovery unit varies
with trife percent sulfur in the crude and the extent of desulfur-
ization-. Typically, 60 percent of the sulfur in the crude
reaches the sulfur recovery plant.
355
-------
^- A Glaus plant operates at about 473^Fiand one to two
atmospheres. About 19 Btu of heat are requirecf~pe"r pound of
sulfur produced. However, about four pounds of steam per pound
of sulfur are produced in a waste heat boiler. ~This steam can
provide from five to thirty percent of the total refining steam
requirements.
The major equipment for a Glaus plant consists of a
combustor and from two to four reactors, depending on the con-
version desired and the acid gas concentration. Each reactor
has an effluent condenser where elemental sulfur is recovered.
Emissions--Process emissions from Glaus plants are
discussed in Section 3.2.1.3 of Appendix E. A.100,000 bpd re-
finery with a one percent sulfur crude and a 95 percent efficient
sulfur plant will produce 5 to 6 tons/day of sulfur emissions.
Possible sulfur emissions are S02, H2S, COS, CS2, and mercaptans.
It is estimated that there are 200 valves, 800 flanges,
9 pump seals, 20 drains, and 4 relief valves on a typical Glaus
unit. These are the sources of the fugitive emissions of various
sulfur compounds from the Glaus unit. However, sulfur recovery
units sampled during this program were found to be leaking hydro-
carbons at rates considerably lower than other refinery process
units.
4.7.6 --• References for Section 4.7
1. American Petroleum Inst., Committee on Refinery Environ-
mental Control. Hydrocarbon Emissions from Refineries.
API Publication No. 928, Washington, D.C.,~1973.
2. "Soudek, M. What Lube Oil Processes to Use. Hydrocarbon
Pro.c., 53(12): 59-66, December 1974.
356
-------
Johirson, Ressell R. and Lawrence E. Swabb^. Jr. The
Importance of Hydrogen in Refinery Operations -Today and
in the future. Eighth World Petroleum Congress, Proc. 4.
67-72, 1971.
Emphasis on H2 Strengthened. Oil and Gas .J., 70(7): 87,
February 14, 1972.
Hydrogen, Steam Reforming. Hydrogen Proc., 52(4): 129,
April 1973.
Schlinger, W. G. and W. L. Slater. Partial Oxidation -
A Minimum Pollution Route for Hydrogen Manufacture. 16(2),
Texaco, Inc., ACS, Div. Petrol." Chem., Montebello, CA,
1971.
Environmental Protection Agency. Development of Petroleum
Plot Plans. EPA-450/3-78-025. Research Triangle Park,
NC, 1978.
357
-------
4.8 ^.. WASTE TREATMENT _ i
*T ' . ^__
4.8.1 Slowdown/Flare Systems
Slowdown/flare systems are items common to all
crude oil refineries. These systems are present for'reasons
including:
• The collection and destruction of small
quantities of waste gas produced during
normal operations,
• The collection and destruction of large
waste gas flows during unit upsets,
shutdowns, fires, or other emergency
situations, and
• The prevention of the above gas streams
from venting directly to the atmosphere
where they might cause safety problems,
contribute to air pollution, or produce
a fire or explosion hazard.
4.8.1.1 Process Description--
A blowdown or pressure-relieving system consists of
relief^yalves, safety valves, manual bypass valves, blowdown
headers, knockout vessels, and holding tanks. A blowdown-
recovery system may also include compressors and vapor surge
vessels. In either case, a flare is used for final disposal of
non-condensible combustible gases.
358
-------
^- An example of a typical refinery blpw^own system is
shown in Figure F4-27. For reasons of safety^arTd -economy,
vessels and equipment relieving to the blowdown system are
usually segregated according to their operating"—pressure. That
is, there is a high-pressure blowdown system for equipment
operating, for example, above 100 psig, and a low-pressure sys-
tem for vessels and equipment operating below 100 psig. In
addition, separate collection systems are provided for
liquids and for a fuel gas purge stream.
The above streams enter the main blowdown drum.
Liquid from this drum is sent to a slops tank while the vapor
is sent to a direct-contact type condenser to permit recovery
of as much hydrocarbon liquid as possible. The resulting light
vapor stream is sent to a flare for disposal by burning.
Butane and propane are usually discharged to a sepa-
rate blowdown drum. This drum is often operated above atmos-
pheric pressure to increase liquid recovery. Liquids from
this drum are recovered while the vapors are combined with the
high pressure blowdown.1
4.8.1.2 Process Technology--
Key elements common to all types of blowdown systems
include pressure relief devices such as safety valves, relief
valves^and rupture discs, and a flare system for the ultimate
disposal of the final waste gas stream. The various types of
pressure relief devices and flare equipment are briefly dis-
cussed below.
359
-------
CO
cr.
o
LOf-PRESSUHE BIO*OG*H
LIQUID BlOnCGXM
FUEL GAS PURGE
HIGH-PHESSURE ElOKCOXN
LIGHT-ENDS
BIOXOOXN DRUH
HA1H
B',0-i(DQXH DRUM
TO FLARE STACK
SCRUBBER
XATER
JLICUIIl-TO-SLCPS.TAh'X
-ZKCS COKO?>1SATE RECOVtRV
Source: Reference 1
Figure FA-27. Typical refinery hlowdown system.
-------
Pressure Relief Devices--Common pressure relief
devices include pressure-relief valves and rupture disks.
Pressure relief valve is a generic term applying to-relief
valves, safety valves and safety-relief valves.
A relief valve is an automatic pressure-relieving
device actuated by the static pressure upstream of the valve.
The valve opens in proportion to the increase in pressure over
the opening pressure. It is used primarily for liquid service.
A safety valve is an automatic pressure-relieving
device actuated by the static pressure upstream of the valve and
characterized by rapid full opening or pop action. It is used
for gas or vapor service. (In the petroleum industry it is
used normally for steam or air.)
A safety-relief valve is an automatic pressure-reliev-
ing device suitable for use as either a safety or relief valve,
depending on application. (In the petroleum industry it is
normally used in gas and vapor service or for liquid.)
A rupture disk consists of a thin metal diaphragm held
between flanges. A rupture disk is sometimes put upstream of a
pressure relief valve to prevent hydrocarbon losses from leaking
pressure relief valve seals.
--• Flares--Flares are used for the disposal of waste
gases by combustion. This includes the combustion of large
surges of gas during an emergency as well as the disposal of
smaller quantities of excess gas produced in more or less con-
stant amounts. There are, in general, three types of flare
systems,: burning pits, elevated flares, and ground level flares.
361
-------
Burning pits are generally reserved~f«r burning large
quantities of gases produced during emergencies. Ordinarily,
the main gas header to the flare system has a water-seal bypass
to the burning pit. Excessive pressure in the header blows the
water seal and permits vapors and gases to vent to the burning
pit.
A typical burning pit is merely an excavation from
4-6 feet deep and from 30-40 feet square, with a burner wall of
brick along one side. The burners are mounted in the wall and
directed at an angle toward the bottom of the pit. Burning pits
are no longer common in U.S. refineries.
An elevated flare consists of a stack, the supporting
structure, the burner tip, and auxiliaries, such as a liquid
trap, seal, pilot burner and igniter. The elevated flare
allows gases to be burned safetly at high elevations and is the
most common type of flare currently in use.
Ground flares are similar to elevated flares. How-
ever, they must be installed in a larger open area than ele-
vated flares for safety and fire protection.2
Smokeless Operation--Smoke is a result of incomplete
combustion and smoke emissions from flares are avoided whenever
possible. Smokeless operation is achieved by following basic
combustion principles:
• Sufficient heat values to maintain critical
combustion temperatures,
• Adequate supply of combustion air, and
362
-------
• Adequate turbulence to promoterniixing
of air and fuel.
Adequate turbulence is usually induced by injecting
steam into the combustion zone. A variety of methods for
mixing steam with air and fuel are available, and all result in
increased velocity and turbulence in the combustion zone. The
use of steam can result in additional benefits:
• Steam reacts with the fuel to form
oxygenated compounds that burn readily
at relatively low temperatures. Water-
gas reactions also occur, with the same
end result.
• Steam reduces the partial pressure of
the fuel and retards polymerization.
Polymerized hydrocarbons are heavier
and more difficult to burn.
Materials other than steam have been used to some
extent to promote smokeless operation; the most common being
air and water. Mechanical air blowing would be ideal for pro-
moting combustion. However, it is generally uneconomical be-
cause the equipment necessary to provide high air flow rate is
quite costly.
Inspiration of air by water has proven successful,
particularly for ground flares. Pumping costs, however, pro-
hibit .the use of water inspiration in elevated "flares. Water
is not quite as effective as steam, however, for controlling
smoke Jszh en high flare rates, olefins, or wet gases are in-
volved.
363
-------
4 . 8 . L. i" Atmospheric Emissions--
Emissions sources from blowdown/flare systems
include: ~~
• Combustion products from flares, and
• Fugitive emissions
Emissions from Flares--In general, emissions of
carbon monoxide and hydrocarbons from flares are higher than
those from process heaters or boilers. Factors which may
account for less effective combustion for flares include:
• Variable firing rates which make control
of steam and combustion air flow rates
difficult.
• Variable heat values for fuel which may
also contain significant quantities of
olefins or aromatics, and
• Relatively low combustion temperatures
with short residence times compared to
process heaters and boilers.
EPA emission factors for smokeless flares are listed
in Table F4-67. The emissions listed here are given as pounds
of pollutant per thousand barrels of refinery capacity.
Fugitive Emissions—Fugitive emissions can occur from
a variety of sources including valves, pumps, Compressors,
flanges and other fittings, pressure relief valves, and drains.
364
-------
TABLE F4-67. EMISSIONS FROM SMOKELESS FLARES
EPA Emission Factors3»*>
(lb/103 Bbl total '
Component refinery capacity)
Particulates Negligible-
SOXC 26.9 , .
CO 4.3
Hydrocarbons 0.8
NOX 18.9
NH3 Negligible
Aldehydes • Negligible
aSource: Reference 3
These emission factors have been given an
emission factor rating of "C."
f\
Varies with fuel sulfur content
Methane content of hydrocarbon emissions was
unavailable
Hydrocarbon emission rates for each source are quite low. How-
ever, total fugitive emissions are significant because of the
large numbers of such sources within the refinery.
The emission factors developed during this program
for each of the above sources are listed in Table F4-68. Ad-
ditional information on these emission factors, including a
complete discussion on their derivation and the confidence in-
tervals for each source category are contained in Appendices B
and C of this report. No attempt has been made here to de-
termine total fugitive emissions from blowdown systems because
of wide variations in the type of waste gas processing em-
ployed^ by refiners.
365
-------
TABLE F4-68.
FUGITIVE EMISSION FACTORS FOR
VARIOUS REFINERY FITTTWGS
Source
Emission Factor
Ib/hr/Source
Valves
Gas-vapor service
Light liquid CVP > O.l psia 0 IOO»F)
Heavy liquid CVP < o.i psia e IOO°F)
Hydrogen service
Open-ended (sample
Pumps seals
Light liquid (VP > 0.1 psia
-------
from coming in contact with process streams, .thts water usually
contains some contaminants. Contaminants enter^he cooling
water via a variety of mechanisms--for example, leaks in heat
exchanger tubes or other process equipment.
Water used directly in processing accounts for a much
smaller portion of the total water requirements than cooling
water. However, this water is almost always laden with con-
taminants due to direct contact with oil. Examples of process
water uses include desalting of crude oil, steam stripping
operations, tank drawoff, and pump gland cooling water.
Although oil free, other water streams such as boiler blowdown
and water from sanitary sources also require treatment for
contaminant removal.
The purpose of the wastewater treating system is to
remove contaminants and provide effluent water of a quality
sufficient to meet state and federal standards. These standards
cover a variety of contaminants and can require the use of
sophisticated treatment techniques.
4.8.2.1 Characteristics of Receiving Waters--
Before beginning a discussion of specific refinery
wastewater treatment.processes, a summary of the effect of
various wastes on water quality is in order. These effects,
summarized from the API manual on the disposal of liquid
refinery wastes, are discussed under the following headings:2
• pH, acidity, and alkalinity,
~"~ • Dissolved oxygen and oxygen demand,
367
-------
^. • Hardness and salinity, _„ ^
-r- .__ ^^
• Toxicity, taste, and odor,
• Color, turbidity, and suspended matter,
• Oil ,
• Temperature.
pH, Acidity, and Alkalinity—The pH value of natural
waters is an approximate measure of acidity or alkalinity in
the water. High or low pH values, caused by the discharge of
refinery waste streams containing strong acids or bases, can
result in a variety of adverse affects.
One important consideration is the bu-ffer effect, or
resistance to change in the pH value upon the addition of
acidic or basic wastes. This buffer effect stems from the
presence of weak acids or bases and various salts. For example,
many natural waters contain carbonic acid (a weak acid formed
by absorption of carbon dioxide), and salts such as sodium or
calcium carbonates and bicarbonates. Moderate amounts of strong
acids or bases can be added to such waters with only small
changes in the pH value. In the absence of such buffering
agents, however, even small additions of strong acid or base
can caus_e a marked change in the pH value. Thus, when eval-
uating" the effects of acidic or basic waste streams, both the
quantity of the waste and the characteristics of the receiving
water must be evaluated.
"~ Low pH values, caused by the additiorr.of untreated
acidie •waste, may result in effects ranging from increased
368
-------
corrosion rates to damage of aquatic life. The^. first effect
of adding strong acid to natural waters is td~cbnvert the
bicarbonates to the salts of the acid added. Carbon dioxide is
liberated during this process. With the continued introduction
of strong acid, the pH value continues to decrease slowly until
a pH of approximately 4.3 is obtained. At this point, all the
bicarbonate has been destroyed and continued introduction of
acid causes the pH to drop sharply.
Acidic waters tend to promote corrosion of steel,
concrete, and.similar structures. The carbon dioxide produced
is itself a corrosive agent and can adversely affect steel,
concrete, and even wooden structures. Highly acidic water is
harmful to many lower forms of life. For example, most bacteria
can not survive at pH values less than 4.5. And, many plankton
are destroyed by a pH value lower than 6.5. These organisms
are essential in the "self purification" processes that normally
oxidize and destroy organic pollution. In addition, the absence
of these organisms tends to break the chain of food supply.
Fish are affected by acidity in a number of ways.
One effect is that acidic water interferes with the ability of
fish to extract dissolved oxygen from the water. An additional
effect is noted in waters where carbon dioxide, produced by the
addition of acidic waste, remains dissolved in the water at high
concentrations. As a result of metabolic processes, fish pro-
duce carbon dioxide which must be eliminated against the partial
pressure of carbon dioxide in the surrounding water. If this
partial pressure becomes too high, as it may from the discharge
of acidic waste, carbon dioxide will not be eliminated fast
enough. The carbon dioxide will accumulate in the tissues,
halting metabolism, and the fish will die. j_
369
-------
^- Highly alkaline wastes, such as caustic from certain
treating processes, can have adverse affects~bn~aquatic life
when discharged to receiving waters. The initial eJffect is a
rise in the pH value of the receiving water. The extent of the
rise for a given quantity of alkaline waste depends on the
buffer action of the natural waters, a situation similar to
that described for the addition of acidic wastes.
Most aquatic life, particularly minute plants, can
not survive when the pH exceeds 9.0. The destruction of minute
plants interrupts the food chain, affecting all forms of aquatic
life. Fish can survive relatively wide ranges in pH. However,
hydroxide ions, present in highly alkaline waters, can attack
the membranes of the gills.
Dissolved Oxygen and Oxygen Demand—Dissolved oxygen
is necessary for the maintenance of fish life.- Generally,
oxygen is present in natural waters at concentrations around
7-9 mg per liter by weight. This oxygen occurs as a result of
absorption of oxygen from the air, and by photosynthesis. When
oxygen is present, an aerobic condition exists and the aquatic
life oxidizes organic matter to carbon dioxide and water. In
the absence of oxygen, an anaerobic, or septic, condition
exists. This results in the breakdown of organic matter to
methane and various sulfur and nitrogen compounds which are
odorous to human beings and toxic to aquatic life. Also,
anaerob_ic waters are corrosive, of poor appearance, and unsuit-
able for many uses.
Many substances present in refinery wastes exert a
demand on the dissolved oxygen content of the water. That is,
they "tlnd to react with dissolved oxygen and remove it from
the water. Oxygen demand can be classified as chemical oxygen
370
-------
demand ^.COD) , biological oxygen demand (BOD) ^and immediate
oxygen "demand (IOD) . "~ ~~ -
Chemical oxygen demand is defined as ±he amount of
oxygen required to oxidize components in the wastewater by
chemical reaction. Several methods have been developed to
measure this characteristic. Biological oxygen demand is
defined as the amount of oxygen required to oxidize the com-
ponents of the wastewater biologically, as determined by an
empirical standard procedure. Immediate oxygen demand refers
to materials which react quickly with dissolved oxygen. This
characteristic is also determined by standard procedures.
Many pollution regulations specify limits on the
discharge of materials which exhibit oxygen demand to prevent
low dissolved oxygen conditions from developing.
Hardness and Salinity — Hardness in fresh water is due
to the presence of dissolved salts. Carbonate or temporary
hardness arises from calcium and magnesium carbonate or
bicarbonates. Noncarbonate or permanent hardness is caused
mainly by calcium and magnesium sulfates and chlorides.
Refinery wastes contain these and other ions in amounts depend-
ing on the quality of the refinery raw water and the .type of
water treatment used.
__. The effect of increasing the hardnes's of fresh
waters is to lessen their value for steam generation and other
industrial uses. The cost of treatment for domestic users is
also increased. .
~ Aquatic life can tolerate increased hardness unless
the concentration of ions is unusually high. Under this
371
-------
circumstance, the water balance within living-.cells can be
disrupted by osmotic effects. ~^
The discharge of salt (sodium chloride) contained in
refinery wastewaters must be carefully monitored. The intro-
duction of large quantities of salt into fresh waters can
damage fish life. Conversely, the introduction of fresh water
into the ocean or its bays can have an adverse effect on marine
life.
Toxicity, Taste, and Odor — Certain refinery wastes
can contain materials toxic to aquatic life. Such materials as
sulfides, ammonia, or phenolics can be directly toxic, through
either internal or external mechanisms, or toxic by indirect
means. The latter group would include the toxic effects of
acids, bases, excessive dissolved solids, and materials which
exhibit oxygen demand.
The toxicity of a particular compound is usually a
function of concentration and time of exposure. A compound
which is harmless at low concentrations may become lethal at
higher concentrations. Some compounds may accumulate in the
tissues and prove lethal at a later time or cause changes in
metabolism, reproduction, or other vital functions. Other
factors affecting toxicity include the characteristics of the
water and the tolerance or immunity developed by various forms
of aquatic life.
Materials such as hydrocarbons, sulfur compounds,
phenolics, and nitrogen compounds can impart ob-jectionable
taste and odor to potable waters, even at low concentration
levels ._ . —,
372
-------
^- . Complete removal of tastes and odorj; produced by some
compounds is extremely difficult and treatment" Tran- even accen-
tuate the problem. For example, if chlorine is used to steri-
lize water which contains phenolics, chlorophen-elics may be
formed. The taste and odor of chlorophenols i's even more of-
fensive than phenolics alone.
Color, Turbidity, and Suspended Matter—The color,
turbidity, and suspended matter contents of water are properties
which may be significantly affected by refinery wastes. The
presence of certain solutes and suspended matter can cause
coloration of the water. Strongly colored water can limit the
penetration of light to the first few feet below the surface.
Thus, photosynthesis, which is an important source of dissolved
oxygen, is reduced.
Turbidity is caused by finely divided particles
which form colloidal suspensions that do not settle. Turbidity
adversely affects recreational uses and the aesthetic qualities
of the receiving water.
As coarser suspended matter settles, it can drag
small plants and animals with it causing the formation of a
sludgy mat on the bottom of the lake or stream. This mat con-
tains organic matter and bacteria which create.a. heavy oxygen
demand.
There are a variety of sources of suspended solids in
refinery wastewaters. These include finely divided carbon,
coke, catalyst fines, clay, silt, cinders, emul_sions, and
flocculent matter.
373
-------
^- Oil--0il discharged with refinery wastewaters is
nearly^always objectionable from a recreationa"f~and aesthetic
viewpoint. Even low concentrations of oil can cause unpleasant
taste and odor problems. Floating oil can coat- the feathers of
water fowl, the gills of fish, and adhere to the surface of
smaller organisms. In addition, heavy oil films interfere with
aeration and photosynthesis. Water soluble components can
exert toxic action on fish and other organisms.
Temperature--The discharge of warm wastewater causes
the temperature of the receiving waters to increase. An
increase in temperature causes the solubility of oxygen to
decrease. Also, reaction rates increase resulting in increased
oxygen consumption. Metabolic functions such as respiration
and food consumption are accelerated. In addition, the toxic
effect of some pollutants increases.
Aquatic organisms are sensitive to temperature and can
exist only within a limited temperature range. Changes in
water temperature can drastically affect the biological makeup
of the receiving waters.
4.8.2.2 Wastewater Treatment Technology--
A variety of potential pollutants is.contained in
refinery waste streams. As indicated in Table F4-69, nearly
every _operation produces waste materials in varying amounts
and forms. Some of these materials are:
• Oil
"^ • Dissolved solids
374
-------
TABLE F4-69
TYPICAL WASTE STREAM CHARACTERISTICS OF
VARIOUS PROCESS OPERATIONS
..a " -M ' ..''A. •
Process Unit
Crude Desalting
Crude Fractlonation
Catalytic Cracking
Thp.rmal Cracking
Hydrocracking
Hydrotreating
OJ
m Delayed Coking
Reforming
Sour Condensates
Alkylation
Flow
Gal/Bbl
2.1
26.0
15.0
2.0
2.0
1.0
1.0
6.0
3.0
60.0
PH
6.7-9.1
8.6
8.3-9.7
6.4
7.3
9.0
8.8-9.1
7.6
4.5-9.5
8.1-12.
BOD
Ib/Bbl
.003
.0002
.015
.001
.002
.010
-
.100
.001
COD
Ib/Bbl
.032
.005
.018
.003
.045
.050
.032
.040
.200
.010
Oil
Ib/Bbl
.012
.017
.100
.001
-
.006
.050
.100
-
H2S NHn
Ib/Bbl Ib/Bbl
.008 .009
.001
.036 .040
.001
.002
.002 .030
.001
1.00 0.75
.010 0
IDS
Ib/Bbl1
.250
.035
.090
-
.002
.035
.030
.125
-
.300
Source: Reference 4 '
-------
. ^.- • Suspended solids _ ~.
• Hydrogen sulfide
• Acids
• Bases
• Soluble organics
• Phenolics
0 Metallic ions
• Ammonia
These materials may be contained in a number of waste
streams including:5
• Oily wastewaters
• Chemical wastewaters
• Boiler blowdown
• Cooling water
• Cooling tower blowdown
• Stripping steam condensates
• Spent catalysts ^
376
-------
• Spent regenerants
• Storm and surface runoff
• Sanitary sewage
The methods used in refinery wastewater treatment can
be classified as follows:
• In-plant or pretreatment
• End of pipe treatment
In-plant or ore-treatment involves the use of procedures
which can (1) reduce the amount of pollutants sent to the waste-
water system, (2) reduce the amount of water discharged to the
receiving waters, and (3) make subsequent end of pipe treatment
more efficient.
End of plant treatment processes can be further
classified as either primary, intermediate, secondary, or
tertiary processes, depending on their function in the treatment
scheme.
In-plant or Pretreatment Technology—A variety of
procedures are available which can reduce the waste load charged
to wate^_treatment facilities. One such procedure is the use of
sour water stripping. Certain water streams, particularly sour
water condensate from stripping operations, contain high concen-
trations of hydrogen sulfide and ammonia. These, components,
which are toxic, odorous, and exhibit high oxygen demand, result
from desulfurization, denitrif ication, and hydro.tr eat ing opera-
tions .--''• • •
377
-------
Hydrogen sulfide and ammonia are often removed via
steam "stripping before this water is sent to^tfie wastewater
system. Processes are available which can recover .essentially
pure sulfur and ammonia. —
Some sour water condensates, particularly those from
cracking or coking operations, contain high concentrations of
water soluble phenol. High concentrations or surges of phenol
can kill bacteria used in subsequent water treatment processes.
To avoid this problem, these condensates are often used as raw
desalter water. A portion of the phenol in the wastewater is
absorbed by the crude oil thereby reducing the amount of phenol
reaching the wastewater system.1*
Other methods of reducing wasteloads to the wastewater
system include limiting waste production at the source. Tech-
niques which limit contact between water and oil such as replac-
ing barometric condensers with surface condensers have been used.
Additional techniques include the installation of initial
treatment facilities for removal of a particular pollutant at
the pollution source.9
Another technique for improving wastewater quality is
to reduce the amount of water which requires treating. A con-
siderable amount of effort has been directed toward developing
methods to recirculate certain water streams rather than dis-
charging them to receiving waters. These techniques include the
installation of air coolers and cooling towers, the elimination
of once-through cooling, chemical treatment to prevent scaling,
corrosion, and biological (algae growth), and a. variety of
others^ These procedures have greatly reduced the quantity of
frestv makeup water required to process each barrel of crude oil.
Addit-ional benefits include a reduction in thermal water pollu-
tion. 5
378
-------
^- There are a number of pretreatment.j:ephniques
available to help improve the efficiency of encf~of pipe water
treatment processes. The most important technique is stream
segregation. It has long been recognized that^water treatment
operations could be made more efficient by treating certain
streams separately from the others. An example of such a stream
segregation scheme is given in Figure F4-28.
By segregating waters which require only minimal treat-
ment from those which require substantial treatment, the size of
treatment facilities can be reduced. Most newer refineries use
a segregated sewer system similar to that of Figure FA-28.
Older refineries, however, were not built with such systems and
the cost of conversion can be prohibitive. Most older refine-
ries do, however, segregate sanitary wastes.
Additional pretreatment schemes include pre-aeration
of the wastewater to meet immediate oxygen demand and equaliza-
tion or surge ponds to smooth variations in water or waste flow.
End of Pipe Treatment—End of pipe treatment consists
of processes designed to produce an effluent water stream of
sufficient quality to meet state and federal pollution regula-
tions. In general, each refinery will have its own particular
water processing scheme based on the type of refinery, the water
use pattern, and the final effluent purity requirements.
The processes used to meet water quality standards are
often classified as shown in Table F4-70. The utilization of
these processes by refiners is indicated in Table F4-71.
379
-------
;TYPE OF ,.,,
WASTE WATER' l\
SOURCES OF WASTES
COLLECTION
5V5TEM
OIL-FREE
OIL* COOLING
OJ
00
o
PROCESS
OIL-FREE STORM WATCR
ONCE- THROUGH COOLING WATER
(C j AWO t.lCHTFR SFRVICE )
COOLING rOwgn SLOWDOWN
(c^ AMD LICHTER SERVICE)
STEAM TURHINE coNOENSEf'-"ATER
AIH COMOITlONfMG COOLING WATE^
HOOF UHAiHAOE
BO'LF.P HLOWOOWN
WATER rRtATMENr PLAMI FILIETI BACKWASH
AMD ION E XCHANGC RECENERATIOfg WASTES
LIQUID TROW WATER ."SOFTENER
SLUDGE DEWATEniNf,
TREATMENT
OR
DISPOSITION
J;SPECIAL-PUTIPOSE;;
SEPAHATOR II
IS
ONCE -THROUGH COOL'WC WATgR
(C» AMD HFAVIER SEfvlCE )
COOL'NG TOwf^R BLOWOOWN
(C« ANO MCAWlcr! SERVICE)
OILYSTOnu WATCH - UNCONTROLLED
OILV5TORM WATER -CONTROLLED
OESAI.TER WATER
1AMK ORAWOPf 5
coMnEMSArs r-noM STRIPPIMC
I'UWPGLANO COOLING WATER
0 4f?OM£ TftIC CONOGM.SER V/AT g
rfifcATiNG PLANT V/ASH WATE^
LOCKED
OILY COOLl'JG WATER
DM5 AND L AVATOO1C3
uGHOur THE PLANT
SANITARY SEWERS
on
riEFWiT I
SE*MCf. fr1
PLAHT j
MUNICIPAL
SEWAGE
Source: Reference 2
Figure F4-28. Example of refinery stream segregation,
-------
TABLE F4-70,
CLASSIFICATION OF REFINERY-WWASTEWATER
TREATMENT PROCESSES —-=-... .
Treatment
Obj ectives
Processes
Pre- or Inplant
Treatment
Primary Treatment
Intermediate Treatment
Secondary Treatment
Tertiary Treatment
Removal of Phenolics, S"
NH3, RSH, Acid Sludge,
Oil Water Reuse and
Equalization
Free Oil and Suspended
Solids Removal
Emulsified Oil, Suspended
Solids, and Colloidal
Solids Removal
Dissolved Organics
Removal Reduction in
BOD and COD
Variable Objectives
Unit Separators
Sfeeam Stripping
Air Oxidation
Neutralization
Surge Ponds
API Separators
Parallel Plate
Separators
Dissolved Air
Flotation
Coagulation-Flotation
Coagulation-Precipita-
.tion
Filtration
Equalization
Activated Sludge
Trickling Filters
Aerated Lagoons
Oxidation Ponds
Rotating Biological
Discs
Filtration
Air Flotation
Coagulation
Activated Carbon
TABLE F4-71. REFINERY UTILIZATION OF WASTEWATER
^ TREATMENT PROCESSES
Primary treatment only
Intermediate treatment
Biologi'ca-1 treatment
• " ~"*^ '
Crude Capacity (BPCD)
3446000
1171000
4433000
9050000
% of US Capacity
33.0
11.2
42.4
r 86.6
Source: Reference 8
381
-------
Primary treatment - Primary treatment ^processes include
API separators and parallel plate separators. The
purpose of these units is to remove floating oil and
suspended solids from the wastewater.T All refineries
use some sort of primary separation device and it is
often the only treatment required.
A. API Separators—API separators have been used in
petroleum refining operations for over 30 years.
Oily water enters one end of a rectangular chan-
nel, flows through the length of the channel, and
discharges at the other end. A sufficient resi-
dence time is provided to allow oil droplets to
float and coalesce at the surface of the waste-
water. An oil skimmer is provided near the end
of the separator to collect floating oil.
Floating oil is advanced toward the skimmer by an
oil and sludge moving device. These devices con-
sist of a series of moving flights which span the
width of the separator. As the flights move over
the surface of the separator, floating oil is
advanced toward the skimmer. The flights return
to the inlet of the separator on the bottom of
the channel. Solids which have settled out of
the water are thus scraped along the channel
bottom to a sludge collecting hopper. Both the
skimmed oil and the settled sludge are collected
and treated for disposal or reuse. The water
stream is either discharged to receiving waters
or routed to additional treating facilities. A
schematic diagram of a typical AP-JE separator is
given in Figure F4-29.
382
-------
00
eovf» F0»fe
" oeimeo
-
r_l
?*j
fc
fx
ik<
~A
><
X
W
(x
l)(J-< SLUDCE PUuP
/
:£M>;W-
r^fi^=
IJ
PLAN
O'L-OC ItN
ftA> H t
* r p-~
—L p~^.
- ^-J"
SECTION A-A
Source: Reference 2
Figure F4-29. Schematic diagram of a typical API separator.
-------
B. Parallel-Plate Separators--The--p^rallel plate
separator is a relatively recent development
which functions by reducing the distance oil
droplets must travel before being~collected and
coalesced. The separator consists of a number of
parallel plates set at an angle and approximately
A inches apart. The oil droplets coalesce on the
underside of the plates and creep up to the
surface of the water. Conversely, solids accumu-
late on the top of the plates and travel to the
bottom of the separator. Parallel plate separa-
tors have the advantage of occupying less area
than API separators.10 Typical efficiencies for
various oil-water separators are.given in Table
F4-72.
Intermediate treatment - API and parallel plate
separators are unable to remove materials which
neither float nor settle within the residence time
provided. Thus, materials such as emulsions and sus-
pended or colloidal solids must be removed by other
methods.
Biological and physical-chemical treatment processes
require that the influent to the treatment system be
of some minimum quality. For this reason, additional
oil and suspended solids removal may .be required after
gravity separation. Such treatment can be provided by
processes such as:
• Dissolved air flotation,
• Induced air flotation,
384
-------
TABLE F4-72. TYPICAL EFFICIENCIES OF OIL SEPARATION UNITS
00
• . oil
Influent
(rtig/D
300
220
108
108
98
100
42
2,000
1,250
1,400
Content .
.1 Effluent
(rag/1)
40
49
20
50
44
40
20
746
170
270
Oil
Removed
87
78
82
54
55
60
52
63
87
81
COD
Removed
Type (%)
Parallel Plate
API 45
Circular
Circular 16
API
API
API
API 22
API
API
ss
Removed
-
-
-
-
-
-
-
33
68
35
:, Reference 10
-------
• Chemical coagulation and sedimentation, and
• Filtration.
Dissolved air flotation (DAF) is used for separating
suspended and colloidal materials from water, includ-
ing suspended solids and insoluble oily wastes. The
DAF unit separates oily wastes and suspended solids
from water by introducing many tiny air bubbles into
the water. These bubbles attach themselves onto oil
globules and suspended solids that are dispersed
through the waste stream. The buoyancy of the resul-
tant oil globule/air bubble complex is substantially
increased. These complexes, therefore, rise through
the wastewater and collect on the water's surface
where they can be removed by surface skimming devices.
In many cases, removal of emulsified oils is enhanced
by adding flocculating agents to the wastewater.
These materials improve oil and solids removal by
improving floe formation which results in more effi-
cient and rapid solids flotation.11
The most common flow scheme for this type of operation
is diagrammed in Figure F4-30. In this example,
wastewater, combined with a flocculation agent, is
charged to a flocculation chamber. From there, the
water-floe mixture flows to the flotation chamber.
Tiny air bubbles are introduced to the flotation
chamber by dissolving air in a recycled portion of
the clarified effluent. This water is mixed with
air at pressures about 40 psig. This"'water is then
sent to the flotation chamber where, under reduced
386
-------
00
WASTE
FLOCCULAriNG
AGENT
OILY SCUM
Source: Reference 8
FLOCCU
LATION
CHAMUtn*
FLOTATION CHAMBER
•
Tv-^
\.
CLARIFIED
EFFLUENT
PRESSUME RETENTION TANK >--\ RECYCLE PUMP
Figure F4-30. Recycle-flow pressurization scheme for a
dissolved air flotation unit.
-------
pressure, the dissolved air emerges-in the form of
small bubbles.8
In addition to removing oil emulsions and suspended
solids, dissolved air flotation can also result in a
reduction in the oxygen demand of the water.
Induced air flotation (IAF) systems operate using
the same principles as the pressurized DAF units.
The gas, however, is self-induced by a rotor-dis-
perser mechanism. The rotor forces, the liquid through
the disperser openings, thereby creating a negative
pressure that pulls the gas downward into the liquid.
IAF oil removal efficiencies are usually comparable
to the DAF systems.12
Coagulation and precipitation have been applied to
refinery wastewaters for removal of suspended and
colloidal pollutants. The conventional coagulation
system utilizes a rapid mix tank followed first by
slow agitation of the mixture in a flocculation cham-
ber to promote floe particle growth and then sedimen-
tation by gravity. The applicability of this process
is determined on a case-by-case basis. The significant
parameters in this process include pH", flocculation
time and chemical selection.7
Filtration is another candidate process for the
removal of oil and solids in refinery wastewater prior
to biological treatment. Several types of filtration
devices have been proposed for removing free and
emulsified oil from refinery wastewat-e-rs. These fil-
ters range from units with simple sand media to those
containing special media which exhibit a specific
388
-------
affinity for oil. The wastewater TSa flow up through
a graded sand media which serves as a filtering and
coalescing section. Thus, even small particles are
separated and retained on the media..- Regeneration of
the filter beds is accomplished by backwashing the
filter pad to remove the solids and remaining oil.
Equalization of wastewaters discharged from refinery
processing units as an intermediate step in the treat-
ment system has been found to greatly improve treat-
ment results. Biological processes as well as
physical-chemical systems operate more efficiently if
the composition and flow of the wastewater feed is
relatively constant. Periodic and unpredictable large
discharges can occur in any refinery. Requirements of
the NPDES require a fairly consistent effluent dis-
charge. : l
Secondary treatment - Petroleum refineries have waste-
water streams containing dissolved organics. Removal
of these organics is required before reuse or discharge
of these streams. Processes for removal of dissolved
organics are generally classified as physical, bio-
logical, or chemical treatment. Wastewaters from the
petroleum refining industry are generally biologically
treated after secondary deoiling. Physical and chemi-
cal treatment are considered advanced treatment pro-
cesses that follow biological treatment.
All of the biological methods for treating organic
wastes in water involve oxidative decomposition by
microorganisms. The biological wastewater treatment
methods used by the refining industry include:
389
-------
• Activated sludge,
• Trickling filters!
• Aerated lagoons,
• Oxidation ponds, and
• Rotating biological discs.
The processes used in each refinery are determined by
the flow and contaminant characteristics of the waste
stream to be treated. As refineries expand or process
configurations change, additional units are built in
series or parallel to accommodate the effluent treat-
ment requirements.
A. Activated Sludge—Activated sludge processes are
used extensively for removing biodegradable
organic contaminants. A high quality effluent
is usually obtainable from a properly designed
and operated activated sludge system.
This process is carried out in a reaction tank
containing a high concentration of microorganisms
Oxygen is supplied to the wastewater in the reac-
tion tank either by mechanical aerators or a
diffused-air system. Microorganisms remove the
organic materials by biochemical synthesis and
oxidation reactions, converting it to COa, water,
and new cell material. The biological solids
must be removed for disposal. "'
390
-------
The main components of the process are the aera-
tion (reaction) vessel and the clarification
tank. Sludge removed from the clarification tank
is recycled to the aeration vessel to maintain
the required concentration of microorganisms. A
portion of the sludge must be discarded and
requires disposal.11
B. Trickling Filters—A trickling filter consists
of a large, open-topped vessel containing a packed
medium that provides a growth site for biological
microorganisms. The soluble organic pollutants
are consumed by the organisms, and converted to
C02, water, and new protoplasm. - In general, the
filter media is corrugated plastic sheets (poly-
styrene and PVC) welded or bonded together in
predetermined configuration. Rock, tile, coal
or slag can also be used; however, problems such
as plugging, resistance to air flow, and low
loadings make these materials less satisfactory.
The air flow through a trickling filter is
generally down and is generally cool.
C. Aerated Lagoons--An aerated lagoon is a lagoon
constructed to a depth of about 10 feet with
oxygen supplied by artificial means. The oxygen
requirements of the system are usually satisfied
with the use of mechanical aeration. This treat-
ment method involves biological oxidation in
which microorganisms convert dissolved or sus-
pended organic contaminants to stable organic
compounds, CO?, and water. This--wastewater
treatment method has been primarily adopted in
391
-------
warm climates and is best used- »s polishing
process after recovery of organics such as phenol.
The common classes of aeration equipment most
often considered are.- diffused_ aeration systems,
submerged turbine aerators, and high- and low-
speed surface aerators. In addition, gravity
aerators and spray aerators are available.
D. Oxidation Ponds--The depth of the ponds is
normally limited to three to four feet to assure
an adequate supply of oxygen so that aerobic
conditions are maintained without mechanical
mixing. Aeration is achieved by oxygen transfer
at the surface and by the photosynthetic action
of algae present in the pond. Microorganisms then
cause aerobic degradation of organic contaminants
in the wastewater.
E. Rotating Biological Discs—A rotating biological
disc system is a mechanical process that brings
wastewater, air, and microorganisms together for
biological oxidation. This process consists of
a series of closely spaced discs (10-12 feet
in diameter) which are mounted on a horizontal
shaft and rotated with about one-half of the
surface immersed in the wastewater. The process
has been used in Europe for several years. The
discs are typically constructed of light-weight
plastic. When the process is ptaced in operation,
the microbes in the wastewater begin to adhere to
the rotating surfaces and grow there until the
entire surface area is covered with a 1/16-1/8
inch layer of biological growth. As the discs
392
-------
rotate, they carry a film of tfas|t:ewater into
the air where it trickles down the surface of
the discs, absorbing oxygen. Upon completion
of a rotation, the aerated and.partially treated
wastewater is mixed with the balance of the
wastewater. This adds to the dissolved oxygen
content and reduces the concentration of organic
matter in the tank. BOD removal and oxidation
of ammonia nitrogen is inversely proportional to
the hydraulic loading on the disc units.11
Tertiary treatment - Some refineries provide additional
treatment downstream of the biological treatment units.
This polishing treatment may be required by changes in
refinery effluent water quality or effluent regulations
which force refiners to install equipment to remove
contaminants remaining after biological treatment.
Facilities commonly utilized by the refining industry
to polish final effluent streams include various units
for the reduction of suspended solids and carbon
adsorption units for organic pollutants removal.
A. Suspended Solids Reduction--Effluent quality can
be improved by further suspended solids reduction
after biological treatment. Suspended solids
removal processes include filtration, air flota-
tion, chemical coagulation with"gravity precipi-
tation, and polishing ponds. Filtration and air
flotation have been applied to biologically
treated refinery wastewaters. The primary
objective of these processes is suspended solids
removal; however, a natural and"corresponding
reduction in organic parameters such as BOD, COD,
TOC, and oil may occur. The magnitude of
393
-------
organic parameter reduction varies among re-
finery effluents.
B. Carbon Adsorption Organics Removal--The activated
carbon adsorption process has no_t been used
widely for refinery wastewater treatment, yet
preliminary pilot work and limited experience
indicate selected applicability. Carbon absorp-
tion systems will remove non-biodegradable and
toxic organics which may be present in the waste-
water after biological treatment,
4.8.2.3 Atmospheric Emissions from Wastewater Treatment Units--
Hydrocarbons are released to the air as fugitive emis-
sions from all of the above named operations. The extent of
these emissions is a function of the amount of hydrocarbon enter-
ing the unit, the volatility of that hydrocarbon, the type of
control built into or onto the unit, and other factors. Thus,
the greatest opportunity for emission should be at the front end
of a wastewater system, i.e., sewers, open ditches, holding ponds
occurring prior to the API separator, and the API separator it-
self. Since the API separator removes most of the hydrocarbons
with the skimmed oil, units following it in the system should
release substantially lower quantities of fugitive hydrocarbons.
^ Typical controls available to the industry include:
operating a sealed sewer system with the hydrocarbon vapors
being flushed to a heater or flare; using fixed or floating
roofs .on the API separator and in some cases venting the hydro-
carbqn^vapors to a heater or flare; eliminating all ditches and
holding-ponds, prior to the API separator; and ttsing covered DAF
units. '.
394
-------
The quantity and character of emissjofis are subject to
between-refinery variations due to the type of crude being pro-
cessed, the process units and product slate of a refinery, the
type of wastewater system, the presence or absence of various
emission controls within the wastewater systemy and the refinery
maintenance practices. The quantity and character.of emissions
also vary within a refinery due to upsets, spills, and turn-
arounds, storm runoff if sewers are not segregated, malfunction-
ing controls in the wastewater system and changes in ambient
temperatures, wind direction and velocity. Several other less. .
common factors also affect emissions such as the presence of
emulsions, malfunctioning of the wastewater unit, improper
bacterial growth, etc.
The current hydrocarbon emission factors listed by the
EPA for API separators are given in Table F4-73. Research is
now underway by EPA to update these figures.
4.8.3 Sludge and Solids Treatment/Disposal
A variety of types of solid waste are generated by
petroleum refineries. Many of the solid wastes contain toxic
hydrocarbons or metallic compounds, which must be handled care-
fully to protect human health and avoid environmental damage.
Refinery solid waste streams may be classified in two
main groups: those wastes that are generated intermittently and
those that are generated continuously.13 Intermittent wastes
include process vessel sludges, scale and other deposits removed
during-turnarounds; storage tank sediments; and~product treat-
ment.wastes, e.g., spent filter clay or catalyst.
: Continuous wastes are those requiring relatively con-
tinuous (at less than two-week intervals) disposal. Typical
395
-------
TABLE-F4-73. API SEPARATOR HYDROCARBON EMISSEDN FACTORS
Emissions
a,b,c
lb/103 Gal - lb/103 Bbl
Wastewater Refinery feed
API separators (uncontrolled)
API separators (controlled by
fixed or floating roof or
vapor recovery system)
5
0.2
200
10
Source: Reference 3
These emission factors have been given an emission factor
rating of "D."
°Less than 1 percent of total emissions are methane.
continuous wastes include coker wastes; spent catalysts and
catalyst fines from fluid catalytic cracking units; spent and
spilled grease and wax wastes from lube oil plants; waste bio-
logical sludges and other oily sludges from wastewater treatment
facilities.
In the past, the majority of refinery solid waste was
either sent to landfills or ponds on-site, or partially de-
watered and hauled to an off-site landfill or dump. Although
landfills are still a common solid waste disposal method,
increasingly stringent environmental regulations have forced
refiners to consider other disposal options, such as incinera-
tion, chemical fixation, or a combination of methods.
The remainder of this section discusses the sources
and characteristics of typical refinery solid wastes, and the
options available for treatment and disposal of these wastes.
396
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4. 8. 3. F Sources and Composition of Ref ineryCS&lid Waste--
Refinery solid waste is generated both from process-
related sources and from effluent treatment units. A list of
wastes from both sources is characterized in Table F4-74. Brief
descriptions of each waste type are given in the following para-
graphs .
Storage tank bottoms accumulate in storage tanks and
are removed when tanks are periodically cleaned. Crude storage.
tank bottoms consist of a mixture of iron rust, clay, sand,
water, and emulsified oil and wax. Product tank bottoms vary
in composition depending on the product stored. Bottoms from
leaded product storage tanks contain small amounts of tetraethyl
or tetramethyl lead, other heavy metals, and phenols.
Neutralized alkylation sludge is produced by both sul-
furic acid- and hydrofluoric acid-based alkylation processes.
The sludge from sulfuric acid processes contains sulfuric acid,
polymerized hydrocarbons, and storage tank scale. Spent hydro-
fluoric acid is neutralized with lime to produce an insoluble
calcium fluoride sludge.
Filter clays are used in removing color bodies, chemi-
cal treatment residues, and trace moisture from gasoline, kero-
sene, jet fuel, light fuel oil, and lube oil streams. Spent
filter~clay forms a sludge or cake which contains traces of oil
and heavy metals.
Catalyst fines are continuously generated by fluid
catalytic cracking units and collected by electrostatic precipi-
tatorsr ; The fines, which are generated continuously, contain
aluminum silicate particles onto which heavy metals (e.g.,
vanadium and nickel) are adsorbed.
397
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i- TABLE F4-74.
SOURCES AND CHARACTERiSSICS
OF REFINERY SOLID WASTES -
Description
Source
Characteristics
Process Solids
Crude oil sediment
and water
Catalyst fines and
spent catalyst
Coke fines
Spent sludges
Spent clay
Crude oil storage,
desalter
Fluid catalytic
crackers, catalytic
reformers, hydro-
treat ing/hydrore fin ing
processes
Coker
HF alkylation units
Lube oil treatment
units
Tank bottoms Storage tanks
Effluent Treatment Solids
API separator API separator
sludge and slop oil
emulsion solids
Chemical precipita- Chemical wastewater
tion sludge treatment processes
Scums or froth
Waste^biological
sludges
Air flotation
Biological treatment
Iron rust, iron sulfides,
clay, sand, water and oil
Inert solids, catalyst
particles (miscellaneous
heavy metals), carbon
Carbon particles,
hydrocarbons
Calcium fluoride,
bauxite
Clay, acid sludges, oil
Oil, water, solids
Oil, sands, and
miscellaneous process
solids
Aluminum or ferric
hydroxides, calcium
carbonate
Oil, solids, flocculants
Water, inerts, biological
solids
Source: "Reference 15
398
-------
Spent catalysts result from a number jbf refinery pro-
cess units which use a fixed-bed catalyst: catalytic reforming,
hydrodesulfurization, hydrotreating, hydrocracking, and others.
Spend catalysts contain a number of heavy metals, plus organic
contaminants absorbed from process feedstocks. -
API separator sludge and slop oil emulsion solids are
non-recycled components from API separators. Slop oil emulsion
solids contain water, oil, and suspended particulate matter.
Separator sludge contains a mixture of all sewered wastes, in-.
eluding tank bottoms, desalter waste and oil. The sludge usually
contains high concentrations of phenols and heavy metals.
Chemical precipitation sludge is generated at some re-
fineries which chemically treat aqueous waste streams. Chemical
coagulants are added to the streams to form a gelatinous precipi-
tate in which oil and solids are trapped.
Waste biological sludge contains oil, suspended solids,
and organic cell material generated in secondary biological
wastewater treatment. Some heavy metals may also be present.
4.8.3.2 Treatment and Disposal of Refinery Solid Wastes--
Both process and effluent treatment solid wastes have
historically been sent directly to landfills or open pits for
dispo&al. Oily wastes, although sometimes incinerated, have
usually been sent to an oily waste disposal pit.
- However, promulgation of more stringent solid waste
disposal regulations has forced refiners to consider methods of
waste-disposal which minimize or prevent discharge of waste-
contaminated materials into the environment. Landfills, although
399
-------
still a^yiable means of waste disposal, are beiiig redesigned to
contain wastes so that no hazardous or toxic materials enter the
air, ground, or surface waters. Other treatment methods, such
as incineration or chemical fixation, are being considered for
use in conjunction with landfilling. Some states permit land-
farming, in which small amounts of oily sludge are distributed
in topsoil. These methods are described in the remainder of
this section.
Landfilling is the most often-used solid waste dis-
posal method. However, landfill design criteria are currently
undergoing change. Regulations proposed under the Resource
Conservation and Recovery Act (1978) require that landfills be
designed and constructed so that no direct contact exists be-
tween the landfill and surface or groundwater.l5 These regula-
tions generally prohibit landfilling of ignitable, reactive, or
volatile wastes, and bulk liquids, semi-solids," and sludges.
Landfills must be lined with either a natural impermeable mate-
rial or one of several artificial liner systems.15 The proposed
regulations require that groundwater and leachate monitoring
systems be maintained around the landfill.15
Landfilling is currently the typical disposal method
for most of the solid wastes listed in Table F4-74. However,
existing landfills are probably inadequate for disposal of most
oily refinery waste, since leaching of toxic organics and heavy
metals^ts a strong possibility.
Landfarming is used in some areas for the disposal of
refinery solid wastes. This method uses soil bacteria to bio-
degrade organic materials in solid wastes. Sludge biodegrada-
is dependent on its oil content and depth of plowing.16
400
-------
~~ No long-term data exist as to the adequacy of land-
farming as a refinery sludge disposal method. Little is known
about the nature of its degradation products or .about the per-
colation and leaching of heavy metals or toxic organic compounds
into ground or surface waters.16
Incineration is a relatively expensive means of solid
waste treatment. Pollution controls may be required on some
incinerators, and additional fuel may be necessary to incinerate
low heat content materials. Additionally, a pretreatment de-
watering step is usually required to decrease the water content
of the solid waste.17
Several types of incinerators may be .used to combust
solid waste: box furnaces, vortex incinerators, rotating hearth
incinerators, rotary kiln incinerators, and fluidized bed incin-
erators.17'18 Supplemental fuel requirements for each type of
incinerator have not been reported in the literature. However,
one source reported that sludge having a heating value of 57,000
Btu/gallon is adequate for supporting combustion in a fluidized
bed incinerator.}6
Although incineration reduces sludge volume, inciner-
ated ash must still be disposed. Disposal is usually to a
landfill, where design and operating constraints discussed pre-
viously for raw sludge may also apply.
Chemical fixation is a solid waste treatment method
in which polyvalent metal ions and waste hydrocarbons are chem-
ically_bound in a cross-linked, polymeric matrix.22 In the
second phase of the process, the polymeric material reacts with
soluble,.silicates in the reactor to form an insoluble solid.
The resulting solid is usually landfilled.
401
-------
--" Results of tests conducted under a^ysriety of simu-
lated landfill conditions have indicated that little leaching
of heavy metals and organic compounds results from chemically -
fixed waste.19
4.8.4 References for Section 4.8
1. Air Pollution Engineering Manual. John A. Danielson, comp
and ed. AP-40. Cincinnati, Ohio. PHS, 1967, pp 565-606.
2. American Petroleum Institute. Manual on Disposal of Re-
finery Wastes. Washington, B.C., 1969.
3. Environmental Protection Agency. Compilation of Air Pol-
lutant Emission Factors. Third Edition, Supplement No. 8.
AP-42. Research Triangle Park, North Carolina, May 1978.
4. Willenbrink, Ron. Wastewater Reuse and In-Plant Treatment.
AIChE Symp. Ser., 70(136): 671-674, 1974.
5. . Kilpert, Richard. Petroleum Refinery Effluent Quality
Control. Regional Workshop on Water Resources, Environment
and National Development, Singapore, March 13-17, 1972.
6. Roffman, Haia K. , and Amiram Roffman. Water That Cools But
Does Not Pollute. Chem. Eng., 83(13): 167-174, June 21,
1976.
7. Grutsch, J. F., and R. C. Mallatt. Optimize the Effluent
System. Hydro. Proc. 55(5): 221-230, May 1976.
8. Volesky, B., and S. Agathos. Oil Removal from Refinery
Wastes by Air Flotation. Water Pollution Research in
Canada, 9:328, 1974.
9. Grutsch, J. F., and R. C. Mallatt. Optimize the Effluent
System. Hydro. Proc., 55(4): 207-209, April 1976.
10. Miranda, Julio G. Designing Parallel-Plates Separators.
Chem. Eng. 84(3): 105-107, January 31, 197_7.
11. Radian Corporation. Atmospheric Hydrocarbon Emissions
--from Petroleum Refinery Wastewater Systems. Technical
. Pjtoposal. Austin, Texas, June 18, 1979. "-(Unpublished)
402
-------
12. F£rd, D. L. , and R. L. Elton. Removal of |)il and Grease
from Industrial Wastewaters. Chemical Engineering,
Deskbook Issue, October 17, 1977.
13. Sittig, Marshall. Petroleum Refining Industry Energy
Saving and Environmental Control. Park Ridge, New Jersey,
Noyes Data Corporation, 1978.
14. U.S. Environmental Protection Agency, Office of Research
and Development. Environmental Considerations of Selected
Energy Conserving Manufacturing Process Options, IV,
Petroleum Refining Industry Report. EPA-600/7-76-034d.
Washington, D.C., December 1976.
15. Dewel, William A., Jr. Solid Waste Disposal: Landfilling.
Chem. Eng. 86(14): 77-86, July 2, 1976.
16. Rosenberg, D. G., et al. Assessment of Hazardous Waste
Practices in the Petroleum Refining Industry. PB-259, 097.
Springfield, Virginia, National Technical Information
Service, June 1976.
17. Stichting Concawe Special Task Force. Incineration Dis-
posals of Refinery Wastes. Oil and Gas J. 73(46): 60-63,
November 17, 1975.
18. Mallatt, R. C., et al. Amoco Launches Two-Pronged Attack
on Pollution. Oil and Gas J. 68(26): 100-114, June 29,
1970.
19. Wisniewski, Ralph. Process Converts Sludge to Landfill.
Oil and Gas J. 73(11): 133-35, March 17, 1975.
403
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SECTION 5
CONVERSION FACTORS
To Convert From
Btu
bbl
gal
ton
Ibs
cm
ft3
psi
g/gal
Btu/bbl
kWh/bbl
Ib/bbl
lb/106 Btu
grain/ ft3
gal/106ft3
gpm
lb/1000 gal
To
kcal
a
a
kg
kg
in
m3
kg/ cm2
g/i
kcal/i
kWh/S,
kg/£
• g/Mcal
g/m3
£/106m3
m3/hr
mg/£
.-Multiply By
0.252
159.0
3.785
907.2
0.454
0.394
0.0283
14.223
0.264
0.0016
0.0063
0.0285
18.0
2:29
133.7
0.227
119.8
404
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TECHNICAL REPORT DATA
(Please read Instructions on the reverse before completing)
\. REPORT NO.
EPA-600/2-80-075e
A. TITLE ANDSUBTITLE
Assessment of Atmospheric Emissions from
Petroleum Refining: Volume 5. Appendix F
5. REPORT DATE
April 1980
6. PERFORMING ORGANIZATION CODE
7 AUTHORIS)
i PERFORMING ORGANIZATION REPORT NO
R.G. Wetherold
9. PERFORMING ORGANIZATION NAME AND ADDRESS
Radian Corporation
P.O. Box 9948
Austin, Texas 78766
10. PROGRAM ELEMENT NO.
1AB604
11. CONTRACT/GRANT NO.
68-02-2147, Exhibit B
12. SPONSORING AGENCY NAME AND ADDRESS
EPA, Office of Research and Development
Industrial Environmental Research Laboratory
Research Triangle Park, NC 27711
13. TYPE OF REPORT AND PERIOD COVERED
Final; 3/76-6/79
14. SPONSORING AGENCY CODE
EPA/600/13
is. SUPPLEMENTARY NOTES lERL-RTP project officer is Bruce A. Tichenor, Mail Drop 62,
919/541-2547.
is. ABSTRACTrpne report gives results of a 3-year program to assess the environmental
impact of petroleum refining atmospheric emissions. This volume gives a detailed
characterization of petroleum refinery technology, prepared as part of the program.
Petroleum refineries in the U.S. are listed, characterized, and classified according
to their types and complexities. It describes four types or sets of refinery models
which could be used to simulate the entire refinery industry. It gives the character-
istics of crude oils, other raw materials, and intermediate and final products. A
major portion of this volume is a detailed description of current refinery process
technology and auxiliary operations. It describes the purpose, operation, energy
needs , and utility requirements. It includes simple flow diagrams for most pro-
cesses. It gives a detailed estimate of the number of fugitive emission sources for
each process type. It also includes detailed estimates of process and fugitive non-
methane hydrocarbon emissions for major refinery process units.
17.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b. IDENTIFIERS/OPEN ENDED TERMS
c. COSATI Field/Group
Pollution
Petroleum Refining
Assessments
Properties
Hydrocarbons.
Pollution Control
Stationary Sources
Technology
Characterization
Nonmethane Hydro-
carbons
13B
13H
14B
07C
13. DISTRIBUTION STATEMENT
Release to Public
19. SECURITY CLASS 'ThisReport)
Unclassified
21. NO. OF PAGES
20. SECURITY CLASS (This page)
Unclassified
22. PRICE
EPA Form 2220-1 (9-73)
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