EPA-450/3-75-028
March 1974
EMISSIONS FROM PROCESSES
PRODUCING CLEAN FUELS
U.S. ENVIRONMENTAL PROTECTION AGENCY
Office of Air and Waste Management
Office of Air Quality Planning and Standards
Research Triangle Park, North Carolina 27711
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EPA-450/3-75-028
EMISSIONS FROM PROCESS!
PRODUCING CLEAN FUELS
by
F. Glazer, A. Hershaft, and R. Shaw
Booz-Allen and Hamilton Inc.
4733 Bethesda Avenue
Bethesda, Maryland 20014
Contract No. 68-02-1358
EPA Project Officer: William Herring
Prepared for
ENVIRONMENTAL PROTECTION AGENCY
Office of Air and Waste Management
Office of Air Quality Planning and Standards
Research Triangle Park, North Carolina 27711
March 1974
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NOTICE
This document includes technical information and recommendations submitted
by the Contractor to the United States Environmental Protection Agency (EPA)
regarding the subject industry. The report, including the recommendations,
will be undergoing extensive review by EPA, Federal and State agencies,
public interest organizations, and other interested groups and persons. The
report and in particular the Contractor's recommended standards of perfor-
mance are subject to change in any and all respects. Mention of company
or product names is not to be considered as an endorsement by the Environ-
mental Protection Agency.
The regulations to be published by EPA under Section 111 of the Clean Air
Act of 1970 will be based to a large extent on the report and the comments
received on it. However, EPA will also consider additional pertinent
technical and economic information which is developed in the course of
review of this report by the public and within EPA. Upon completion
of the review process, and prior to final promulgation of regulations, an
EPA report will be issued setting forth EPA's conclusions concerning the
subject industry and standards of performance for new stationary sources
applicable to such industry. Judgments necessary to promulgation of
regulations under Section 111 of the Act, of course, remairt the responsi-
bility of EPA. Subject to these limitations, EPA is making this report
available in order to encourage the widest possible participation of interested
persons in the decision-making process at the earliest possible time.
This report shall have standing in any EPA proceeding or court proceeding
only to the extent that it represents the views of the Contractor who studied
the subject industry and prepared the information and recommendations.
It cannot be cited, referenced, or represented in any respect in any such
proceedings as a statement of EPA's views regarding the subject industry.
This report is issued by the Environmental Protection Agency to report
technical data of interest to a limited number of readers. Copies are avail-
able free of charge to Federal employees, current contractors and grantees,
and nonprofit organizations - as supplies permit - from the Air Pollution
Technical Information Center, Environmental Protection Agency, Research
Triangle Park, North Carolina 27711; or, for a fee, from the National
Technical Information Service, 5285 Port Royal Road, Springfield,
Virginia 22161.
Publication No. EPA-450/3-75-028
11
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TABLE OF CONTENTS
Page
Number
SECTION 1
I. INTRODUCTION 1-1
II. SUMMARY OF ANALYSIS METHODOLOGY H-l
III. DISCUSSION OF SELECTED TOPICS COMMON
TO ALL PROCESS DESCRIPTIONS HI-1
SECTION 2
IV. THE SYNTHANE PROCESS lv~1
V. THE CO ACCEPTOR PROCESS V-l
VI. THE LURGI PROCESS VI-1
VII. THE HYGAS PROCESS VII-1
VIII. THE COED PROCESS VIII-1
DC. THE SOLVENT REFINED COAL (SRC) PROCESS IX-1
X. THE TEXACO PARTIAL OXIDATION PROCESS X-l
-111-
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Page
Number
XI. THE DESULFURIZATION OF CRUDE OIL XI-1
XII. THE GAS COMBUSTION RETORT PROCESS XII-1
XIII. THE LURGI PROCESS XIII-1
XIV. THE U-GAS PROCESS ; XIV-1
XV. THE KOPPERS-TOTZEK PROCESS XV-1
XVI. BIBLIOGRAPHY XVI-1
-iv-
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INDEX OF FIGURES
Page
Number
III-l Two-Stage Acid-Gas Treatment With
Sulfur Recovery as Applied in the Synthane
Process 111-56
III-2 Two-Stage Acid-Gas Treatment as Applied
in the Lurgi Process 111-57
III-3 Selective Acid-Gas Treatment with Glaus-
Type Sulfur Recovery As Applied in the
Texaco and Hygas Processes 111-59
III-4 Format for Calculating Incremental
Capital Investment 111-62
III-5 Format for Calculating Incremental Annual
Operating Cost IH-63
III-6 Derivation of DCF Equation for Incremental
Price of Clean Fuel 111-67
III-7 Derivation of Utility Financing Equation For
Incremental Cost of Fuel 111-68
IV-1 The Synthane Process IV-3
IV-2 Incremental Cost of Gas IV-21
V-l CO Acceptor Process Flow Sheet V-3
VI-1 The Lurgi Process VI-3
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Page
Number
VII-1 The Hygas Process—Montana Subbituminous VII-3
VII-2 The Hygas Process— Illinous No. 6 Coal VIM
VIII-1 The COED Process —Utah A Seam Coal VIII-3
VIII-2 The COED Process—Illinois No. 6 Coal VIII4
IX-1 The Solvent Refined Coal Process IX-2
X-l The Texaco Process Synthetic Gas
Production X-3
X-2 The Texaco Process Hydrogen
Production X-4
XI-1, Desulfurization of Crude Oil XI-3
Case 1
XI-1, Desulfurization of Crude Oil XI4
Case 2
XI-1, Desulfurization of Crude Oil XI-5
Case 3
XI-1, Desulfurization of Crude Oil XI-6
Case 3A
XI-1, Desulfurization of Crude Oil XI-7
Case 4
XI-2 Typical Flow Scheme for a Fixed Bed
Hydrodesulfurization Process XI-12
XI-3 Typical Natural Gas Reforming Process XI-14
VI
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Page
Number
XI-4 Percent of Crude Oil Desulfurized for
Reduced Sulfur in Desulfurized Oil,
100, 000 BPD Crude Oil, 1. 7 Percent Sulfur XI-22
XI-5 Capital Investment for Sulfur Reduction for
100, 000 BPD Crude Oil, 1.7 Percent Sulfur XI-32
XI-6 Cost of Sulfur Reduction for 100, 000 BPD
Crude Oil, 1. 7 Percent Sulfur XI-35
XII-1
The Gas Combustion Retort Process
XII-3
XIII-1 The Lurgi Process (Airblown Gasifier)
XIII-3
XIV-1 The U-Gas Process—Combined Cycle
Case
XIV-2 The U-Gas Process—Fuel Gas Case
XIV-2
XIV-4
XV-1 The Koppers-Totzek Process
XV-2
VII
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INDEX OF TABLES
Page
Number
1-1 Summary of Selected Wastewater Treatment,
Sulfur Removal and Sulfur Recovery
Processes 1-6
1-2 Summary of Costs for Pollution Control 1-9
1-3 Synthetic Fuel Plants Recommended for
Project Independence 1-12
II-1 Coal to Gas Conversion II-4
II-2 Conversion of Coal to Coal or Oil II-6
II-3 Conversion of Lightweight Oil to Gas II-8
II-4 Conversion of Middleweight Oil to Gas II-9
II-5 Conversion of Heavyweight Oil (and Tar)
to Gas / II-10
II-6 Oil to Oil Conversion 11-11
II-7 Conversion of Oil-Bearing Minerals
(Tar-Sands, Oil Shale) to Syn-Crude 11-12
II-8 Conversion of Oil Shale to Gas 11-13
II-9 Selection of Clean-Fuel Processes for
Detailed Study 11-15
11-10 .Nature and Emissions of Major Sources
of Pollutants 11-17
11-11 Units Used in Process Display 11-21
Vlll
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Page
Number
III-l Nature and Sources of Major Waste Streams IJI-3
III-2 Composition of Selected Gaseous Waste
Streams (Vol. Percent) III-6
III-3 Composition of Selected Liquid Waste
Streams (mg/1) III-7
III-4 Composition of Selected Solid Waste
Streams (wt. percent) III-8
III-5 Sulfur Distribution Among Outputs of
Clean-Fuel Processes (Weight Percent
of Feedstock) III-l 1
III-6 Distribution of Energy in Clean-Fuel
Products (Percent of Feedstock Energy) III-l3
III-7 Comparison of 250 Billion Btu/day Gas
From Coal with 3000 MW Electricity from
Coal III-14
III-8 Solvent-Based Process 111-39
III-9 Amine Processes 111-43
111-10 Indirect Oxidation Processes 111-45
IV-1 Sulfur Compounds in Scrubber Effluent Gas IV-5
IV-2 Composition of Wastewater From Gasifier IV-5
IV-3 Energy Balance Calculations IV-9
IV-4 Sulfur Balance - Short Tons/Hr (m Tons/Hr) IV-9
IV-5 Nature and Treatment of Major Waste Streams IV-10
IX
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Page
Number
IV-6 Incremental Capital Investment •($ Million) IV-17
IV-7 Incremental Annual Operating Cost ($) IV-18
IV-8 Incremental Cost of Gas Due to Pollution
Control for the Synthane Process IV-20
V-l Compositionsiof Gaseous Streams V-2
V-2 Solids Analysis V-4
V-3 Energy Balance Calculations V-l 1
V-4 Sulfur Balance V-l 2
V-5 Source and Treatment of Major Waste
Streams' V-13
V-6 Incremental Capital Investment ($ Million) V-18
V-7 Incremental Annual Operating Cost ($) V-19
..,.-. ..•••••"
V-8 Incremental Cost of Gas Due to Pollution
Control for the CO Acceptor -Process V-20
£
VI-1 Stream Composition VI4
VI-2 Component Analysis (Dry Coal) VI-5
VI-3 Low-Btu Fuel Gas Composition VI-11
VI-4 Composition and Quantity of Steam and
Power Generation Effluents to the Atmosphere VI-11
VI-5 Energy Balance Calculations VI-13
VI-6 Sulfur Balance VI-14
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Page
Number
VI-7 Nature-and Treatment of Major Waste
Streams VI-15
VI-8 Sulfur Compounds Vented to Atmosphere
from Stretford Units VI-16
VI-9 Incremental Capital Investment ($ Million) VI-21
VI-10 Incremental Annual-Operating Cost"•<$) Vl-22
VI-11 Incremental Cost of Gas VI-23
VII-1 Component; Analysis (Ory Goal') VII-5
•
VII-2 Montana Coal Stream Compositions VII-7
VII-3 Illinois Coal Stream Compositions VII-10
VII-4 Energy Balance VII-18
VII-5 Sulfur Balance VII-18
VII-6 Source,and Treatment of;Major
Waste Streams VII-19
VII-7 Hygas Plant With Montana Subbituminous
Coal Feed, Incremental Capital Costs VII-27
VII-8 Hygas—Illinois No. 6 Coal, Incremental
Capital- Incremental
Annual Operating Costs VII-30
VII-11 Incremental ;Cost of,Gas Duetto .Pollution
Control for the Hygas Process VII-32
XI
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Page
Number
VIII-1 COED Feedstocks and Products VIII-2
VIII-2 Energy Balance Calculations VIII-10
VIII-3 Sulfur Balance VIII-11
VIII-4 Nature and Treatment of Major Waste
Streams VIII-12
VIII-5 Incremental Capital Investment ($ Million) VIII-17
VIII-6 Incremental Annual Operating Cost ($) ' VIII-18
• . ,••••'• , .. / -' ' •• '' . •':
VIII-7 Incremental Cost of Clean Fuel Due to
Pollution Control for the COED Process VIII-20
IX-1 Composition of Gaseous Streams- ' IX-3
IX-2 Composition of Liquid Streams " ' IX-4
IX-3 Composition of Solids Streams IX-5
IX-4 Component Analysis for Coal Feedstock . IX-6
IX-5 Energy Balance Calculations IX-8
IX-6 Sulfur Balance for the SRC Process IX-15
IX-7 Source and Treatment of Major Waste
Streams IX-16
IX-8 Incremental Capital Investment ($ Million) IX-21
IX-9 Incremental Annual Operating Costs ($) IX-22
IX-10 Incremental Cost of Solvent Refined Coal
Product (SRC) Due to Pollution Control IX-24
XII
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Page
Number
X-l Sulfur Balance, Short Tons/Hr (Metric
Tons/Hr) X-5
X-2 Energy Balance X-6
X-3 Nature and Treatment of Major Waste
Streams X-10
X-4 Incremental Capital Investment ($ Million)
For Synthesis Gas Production
(For 6. 2% Sulfur in Residual Oil Feedstock) X-13
X-5 Incremental Capital Investment ($ Million)
For Hydrogen Production
(For 6. 2% Sulfur in Residual Oil Feedstock) X-14
X-6 Incremental Annual Operating Costs of
Synthesis Gas Production X-l5
X-7 Incremental Annual Operating Costs of
Hydrogen Production X-l 6
X-8 Incremental Cost of Synthesis Gas
Production Due to Pollution Control
(6. 2% Sulfur in Residual Oil Feedstock) X-18
X-9 Incremental Cost of Hydrogen Production
Due to Pollution Control
(6. 2% Sulfur in Residual Oil Feedstock) X-18
XI-1 Important Hydrodesulfurization Processes XI-10
XI-2 Hydrogen Plant Requirements and
Production Rates XI-15
XI-3 Steam Balance, Fuel and Power
Requirements XI-20
Xlll
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Page
Number
XI-4 Fuel Gas Production, FOE barrels/day XI-25
XI-5 Expected Water Requirements XI-26
XI-6 Spent Catalyst, Ib/day (kg/day) XI-29
XI-7 Incremental Capital Investment ($ Million) XI-30
XI-8 Incremental Annual Operating Cost ($) XI-33
•XI-9 Summary of HDS Capital Costs ' XI-34
XI-10 -Summary of HDS Operating Costs • XI-34
,XI-11 Incremental Cost of Desulfurizing Crude '
Oil XI-36
XII-1 Oil Shale Component Analysis
XII-2 Energy'Balance Calculations for the Gas
Combustion Retort Process
XII-3 Sulfur Balance for the Gas Combustion
Retort Process
XII-4 Nature and Treatment of Major Waste
Streams . .• -
XII-5 Untreated and Treated Wastewater
XII-6 Incremental Capital Investment
XII-7 Incremental Annual Operating Cost
XII-8 Incremental Cost of Oil Due to Pollution
Control for the Gas Combustion Retort
Process .
XII-4
XII-6
XII-10
XII-11
XII-13
XII-17
XII-18
XII-19
xiv
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Page
Number
XIII-1 Stream Composition
XIII-2 Component Analysis (Dry Coal)
XIII-3 Energy Balance
XIII-4 Sulfur Balance (29. 7 x 109 kcal/day
Plant - Net) (117. 7 x 109 Btu/day)
XIII-5 Wastes, Their Amounts, Sources, and
Treatment for a Coal to Low-Btu Fuel
Gas Plant
XIII-6 Incremental Capital Investment ($1000)
(117. 7 Billion Btu/Day Net Plant)
XIII-7 Incremental Annual Operating Cost ($)
(117. 7 Billion Btu/day Net), 90% Stream
Factor)
XIII-8 Incremental Cost of Gas (117. 7 Billion
Btu/Day (Net) Plant)
XIII-4
XIII-5
XIII-6
XIII-11
XIII-13
XIII-17
XIII-18
XIII-19
XIV-1 Stream Composition —Combined Cycle
Case
XIV-2 Stream Composition — Fuel Gas Case
XIV-3 Component Analysis (Dry Coal)
XIV-4 Energy Balance Calculations, Combined
Cycle Case
XIV-5 Combined Cycle Case, Sulfur Balance
XIV-6 Energy Balance Calculations, Fuel Gas
Case
XIV-5
XIV-6
XIV-7
XIV-13
XIV-14
XIV-16
xv
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Page
Number
XIV-7 Fuel Gas Case, Sulfur Balance XIV-17
XIV-8 Source and Treatment of Major Waste
Streams for U-Gas Process — Combined
Cycle Case (Airblown Gasifier) XIV-18
XIV-9 Source and r.l reatment of Major Waste
Streams for U-Gas Process—Fuel Gas
Case (Oxygen-Blown Gasifier) X1V-22
XIV-10 Incremental Capital Cost for Pollution
Abatement for Combined-Cycle Facility
Using Airblown U-Gas Process XIV-24
XIV-11 Incremental Operating Costs for Pollution
Abatement for Combined-Cycle Facility
Using Airblown U-Gas Process XIV-25
XIV-12 Incremental Capital Cost for Pollution
Abatement for Fuel Gas Plant Based on
U-Gas Process (Oxygen-Blown Gasifier) XIV-26
XIV-13 Incremental Operating Costs for Pollution
Abatement for Fuel Gas Plant Based on
U-Gas Process (Oxygen-Blown Gasifier) XIV-27
XIV-14 Incremental Cost of Gas ($/106 Btu) XIV-28
XV-1 Stream Compositions, Illinois Coal XV-3
XV-2 Stream Compositions, Eastern Coal XV-4
XV-3 Stream Compositions, Western Coal XV-5
XV-4 Coal Composition and Heating Values XV-6
XV-5 Energy Balances for Various Coal Feeds XV-12
xvi
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Page
Number
XV-6 Sulfur Balances XV-13
XV-7 Nature and Treatment of Major Waste
Streams XV-14
XV-8 Koppers Coal Gasification Water Analyses,
Kutahya, Turkey XV-17
XV-9a Incremental Capital Cost for Pollution
Abatement for Koppers-Totzek Process
Using Illinois Coal XV-19
XV-9b Incremental Operating Costs for Pollution
Abatement for Koppers-Totzek Process
Using Illinois Coal XV-20
XV-lOa Incremental Capital Cost for Pollution
Abatement for Koppers-Totzek Process
Using Eastern Coal XV-21
XV-lOb Incremental Operating Costs for Pollution
Abatement for Koppers-Totzek Process
Using Eastern Coal XV-22
XV-lla Incremental Capital Cost for Pollution
Abatement for Koppers-Totzek Process
Using Western Coal XV-23
XV-lib Incremental Operating Costs for Pollution
Abatement for Koppers-Totzek Process
Using Western Coal XV-24
XV-12 .Incremental Cost of Gas XV-25
XVI1
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SECTION 1
I. INTRODUCTION
Processes for producing clean fuel from coal, oil, oil shales,
and tar sands are receiving increased attention today because:
They offer a way to increase supplies of natural gas and
petroleum products at a time when traditional sources
are being depleted
They generate clean fuels which can be burned with mini-
mal release of harmful emissions to the environment.
Many of the clean-fuels processes which have been under develop-
ment during the last decade are now becoming commercially viable
as technical problems are solved, economies of scale are achieved,
and the prices of alternative fuels increase.
From an environmental point of view, these processes are
attractive because the fuel they produce is very low in sulfur content
and contains virtually no other impurities. The impurities which
were in the raw fuel feedstock are removed during the conversion
process, and will be handled in an environmentally acceptable manner
at the plant site. The overall pollution control problem is simplified
by having the pollutants generated at a single processing site rather
than at myriad combustion points. There remain many complex
technical problems to be solved in controlling the emissions from
clean-fuel processes.
1. PURPOSE OF THE STUDY
The objective of the research effort described in this report
was to assess the technical feasibility and to estimate the costs of
installing alternative pollution control systems on clean-fuel processes
which are likely to become commercially viable within the next decade.
The intent was not to establish standards but rather to:
Determine the types and estimated amounts of pollutants
generated in clean-fuel processes
1-1
-------
Determine the degree of availability of processes for mini-
mizing emissions from these processes
Estimate the economic and energy costs for control of
these emissions
Indicate possible means for utilizing the by-products from
these processes in environmentally satisfactory ways.
'
The primary emphasis in the study was placed on the control of sulfur
emissions, although other pollutant streams were considered and
evaluated as well. In order to quantify pollutant streams, it was
necessary to prepare:
Detailed process flow diagrams, including unit processes
for pollution control
Energy and material balances for the process as a whole
and for all major streams.
As a result, the process descriptions which are presented in this
report are in many cases more complete than those appearing in
other sources, particularly since they contain detailed identification
of pollution control streams and processes.
Accomplishing the stated objectives of the research program
was a major undertaking for several reasons:
Many of the processes which may be commercialized
within ten years are still in the early developmental
stages and detailed data are not yet available. Hence,
determining emissions levels from full-scale plants re-
quired extrapolation of preliminary results from pilot or
laboratory tests.
Most process developers have concentrated their efforts
to date on the main process stream producing the clean
fuel and have not yet devoted extensive time to the control
of pollutant side streams. Hence, to analyze the emis-
sions from these processes, it is necessary to postulate
alternative control systems and to estimate emissions
based on test data for these systems operated under •
different conditions.
1-2
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Almost all of the processes can be operated with different
feedstocks with sulfur content ranging from almost zero
percent to over seven percent. The by-product and
pollutant streams will vary considerably in pollutant con-
centration as a result, and the cleanup processes which
are applied will have to be selected in each case to handle
the specific waste streams generated. There are a very
large number of combinations of:
Basic clean-fuel processes
Potential feedstocks and product streams
Alternative cleanup processes,
to be. analyzed. To determine the optimum pollution con-
trol system, from an environmental and cost viewpoint
for each clean-fuel process, would require a commitment
of effort far beyond the scope of this program.
The approach adopted to deal with these problems and to provide
results of value in assessing pollution control alternatives without
undertaking a major engineering design program involved:
Selection of representative processes to study in detail
Discussion with process developers to obtain specific
information not available in the published literature
Application of informed engineering judgment to incor-
porate cleanup processes not specified by the developer
and to extrapolate data and to estimate system performance
characteristics from prototype process results.
The processes selected for detailed study were:
. Synthane (coal to high-Btu gas)
CO2 Acceptor (coal to high-Btu gas)
' Lurgi (coal to high-Btu gas)
Lurgi (coal to low-Btu gas) .
Hygas (coal to high-Btu gas)
U-Gas (coal to low-Btu gas)
Koppers-Totzek (coal to low-Btu gas)
COED (coal to oil)
1-3
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SRC (coal to coal or oil)
Texaco Partial Oxidation (oil to synthesis gas or hydrogen)
HDS (oil to oil)
Gas Combustion Retort (oil shale to oil).
Descriptions of each of these processes are presented in Section 2 of
this report. Although the results given, concerning pollution genera-
tion and control, are more comprehensive than data that are available
elsewhere, it must be emphasized that these results are, in many
ways, premature. Only a few of the above processes have yet been
engineered to include pollution control systems. For others, sig-
nificant modifications to the pollution control processes identified
will doubtless be made before commercial application occurs; this
report anticipates many developments in the control of pollutants from
clean-fuel processes. Although the emission control processes are not
yet proved in the applications presented in this report, this document
nevertheless is a resource which can be used to assess the options that
are currently available and to identify their capabilities and costs.
2. SUMMARY OF MAJOR RESULTS
Most of the clean-fuels processes considered in this report have
not yet been developed to the point of commercialization and, as a re-
sult, the pollutant streams they will generate and the processes which
will be used to control the environmental discharges from these streams
can not yet be precisely defined. However, based on the preliminary
data and engineering estimates which have been developed in this re-
search program, it can be concluded that:
Unit processes apparently exist to provide adequate control
of most waste streams from clean-fuel plants, although
some pollutants remain difficult to treat
The pollutant emissions from clean-fuel processes will,
in general, be relatively low when compared to emissions
resulting from alternative uses of the raw fuel
The total costs of pollutant control (1973 basis) for achiev-
ing low levels of plant emissions, in high-Btu coal gasifi-
cation processes, range from 15 to 30 cents per million Btu
and can be expected to increase in proportion to the severity
of the emission controls required. As a comparison, using
the same basis (1973) the overall cost of manufacturing
high-Btu. gas. from coal is on the order of $1/10° Btu and
1-4
-------
the average regulated wellhead price of natural gas is less
than $0.25/106 Btu. *
The cost of pollution control in the manufacture of low-
Btu gas are more difficult to assess. From one point of
view the overall process of low-Btu gasification followed
by desulfurization may be considered as a pollution con-
trol system because it permits the combustion of coal with
decreased emissions. In that case the overall cost of pol-
lution control is about $0. 65/10 Btu of gas produced.
On the other hand, using the same context as the other
clean fuels processes reported here, the cost of the emis-
sion control facilities alone within these low-Btu coal
gasification processes ranges from $0.10 to $0. 33/106 Btu.
In the manufacture of clean liquid fuels from coal, oil,
shale or high-sulfur crude, the total costs of pollution
control range from about $0.05 to $0. 25/106 Btu. As a
comparison, the controlled price of crude oil at $5.25/ bbl"'
is $0. 90/106 Btu.
Summary results supporting these conclusions are presented in the
discussion which follows and the remainder of this report presents
the detailed data and process analyses from which this summary
information was drawn.
Table 1-1 provides a brief listing of selected processes which
can be applied to the various pollutant or product streams generated
in clean-fuels processes. Emphasis is placed here on the treatment
of streams containing sulfur, although in the process synopses pre-
sented in Section 2 of this report, the removal and recovery of other
by-products and residuals are also discussed. The generic or
licensed cleanup processes included in Table 1-1 are those which
have been applied most frequently to the clean-fuel processes dis-
cussed in this report. Many other alternatives are available and are
discussed later in Chapter III.
The decision to apply a particular control process to a waste
stream is based on the: ,
These prices are based on 1973 data and may be expected to
increase in the future.
1-5
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Table 1-1
Summary of Selected Wastewater Treatment, Sulfur Removal
and Sulfur Recovery Processes
Type of Pollutant or Product
Stream Treatment
Control Process^1*
Remarks
Acid-gas treatment
Wastewater I real men t
• Hot potassium carbonate
• Rcctisol
• Selexol
• Snlfinol '
• ML A (monoethunolaminc)
• DIPA (diisopropanol aminc)
• MDLA
(metliyldiethanolaminc)
An alkaline salt process that removes
H->S and CO-, as well as CS2 and COS.
Sulfur compounds can be reduced to
about 10 ppm.
A solvent based process that removes
CO2, H2S (down to 0.1 ppm) and
the COS"present
A solvent based process that can remove
H2S, COS and CO2
A solvent-amine based system which
removes CO2 down to 10 ppm and H2S
to 1 ppm. CS2 and mercaptans are also
removed.
An amine scrubbing scheme which
removes CO2 and can reduce sulfur
content to 1 ppm. COS, CS2 and
mercaptans are removed from the
feed stream but degrade the MEA.
An amine system which removes t
H-,5 and CO-,
A tertiary amine system that is partially
selective in removal of H2S
• Phenosolvan
• Phosam
• Lime softening, filtration,
zeolite treatment, biological
treatment
Separates NHj, C'Oo and H^S from
liquid waste streams
Removes phenol from wastewater
down to 20 ppm
Recovers NH3
Various treatment processes
presently used
Sulfur Recovery
• Claus
• Stretford
Recovers elemental sulfur from
acid-gas streams containing as little
as 2% H2S. For economical use,
the untreated stream should contain
over 15% H2S. COS and CS2 can
be removed by special treatment.
Tail-gases usually require further treatment.
An indirect oxidation process that is
up to 99% efficient in recovering
elemental sulfur from low H2S content
streams. It is not effective in reducing
COS content-in the feed stream. The
treated stream can contain as little as
I ppm H2S. CO2 is not removed.
1-6
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Table 1-1
(Continued)
Type of Pollutant or Product
Stream Treatment
Control Process' '
Remarks
Allied
Converts SOi to elemental sulfur
Clans tail-gas treatment
• Incineration .
• Beavon
• Scot
• Converts HiS to SOi which may
be further treated.
Can convert most sulfur forms to HiS
which is eliminated in a Strelford unit.
Reduces sulfur content to below 250 ppm,
Similar to Beavon process, but recovers
UTS in DIPA for recycle to a Clans unit.
Stack gas cleanup
Wcllman-Lord
Li me
A sodium-based scrubbing process
which can remove SC>2 down to
less than 250 ppm. Can also be used
for Claus tail-gas treatment following
incineration.
A general type of scrubbing process
currently practical to remove SO-i
from stack gases.
Trace sulfur removal
• Iron oxide, /.inc oxide
or carbon treatment
Also called sulfur guards or guard
drums. Adsorbs traces of sulfur
remaining in the main gas stream
(< 300 ppm sulfur containing
streams).
(I) The selection and order of processes indicated
does not imply a preference based on effectiveness
of treatment or optimized cost.
Composition of that stream
Severity of treatment required.
The waste stream composition in turn depends on the feedstock, the
process conditions, and the product mix desired. Hence, the appli-
cation of the processes indicated in Table 1-1 can only be made after
detailed analysis of the specific clean-fuel operation to be controlled.
Most of the purification processes discussed in this report were
developed for application in systems other than clean-fuel plants.
For example, the processes for removal of sulfurous compounds
such as H2S were developed for gas field applications. Hence, these
purification processes have not yet been specifically engineered for
clean-fuel process applications. The acid-gas cleanup processes
likewise have not been developed specifically for clean-fuels appli-
cations. However, in the context of this document, they have not
1-7
-------
been considered as pollution control systems. In many processes,
acid-gas treatment is required to protect sensitive downstream
catalysts. In the case of low-Btu gases, however, sulfur removal
from the gas stream is considered an emission control process be-
cause sulfur removal is not intrinsic to the process operation.
Nevertheless, the operation of the acid-gas treatment sections of all
the processes is a major factor in the selection and performance of.
downstream pollution control units. The uncertainty of this opera-
tion has been one consideration in the analysis.
The primary pollution control processes considered in this docu-
ment are wastewater treatment and sulfur recovery. The cost of in-
stalling and operating these processes is considered as an additional
cost to the basic process. These costs are listed for each process
as the incremental cost of pollution control (Table 1-2).
The pollution control processes used here have generally not
been applied to clean-fuels processes. These process uncertainties
have been another consideration in this analysis. In addition, it is
likely that new processes will be developed to handle such pollution
control problems as:
Carbonyl sulfide and carbon disulfide generation
Cyanide in wastewater
Thiocyanate formation and its effect on ammonia and
sulfur removal from wash water
which have not yet been fully resolved by clean-fuel developers.
Table 1-2 provides a summary of the principal results obtained
from the analyses of specific clean-fuel processes presented in Sec-
tion II of this report. For each of the 12 processes selected for
study and for each feedstock analyzed, the table gives
The disposition of the sulfur in the feed
The costs for pollution control computed using two
accounting methods
Discounted cash flow
Utility financing.
1-8
-------
Table 1-2
Summary of Costs for Pollution Control
Process
Synlhane
t'O-i Acceptor
Lurgi )Hii:!l-Bui )
Hygas
COED
SRC
Texaco Partial
Oxidation
HDS
Gas Combustion
Retort
Lurgi ( Low-Bin )
U-Gas
Koppers-Totzek
Feedstock
Pittsburgh
Seam Coal
Illinois No. 6
North Dakota
Lignite
Navajo Coal
Montana
Subbituminous
Illinois No. 6
Utah A
Seam Coal
Illinois No. 6
Coal
Kentucky
No. 1 1 Coal
Residual
Fuel Oil
: Persian Gulf
Crude Oil
Colorado
Mahogany Zone
Navajo Coal
Pittsburgh
Seam Coal
Illinois Coal
Eastern Coal
Western Coal
Alternate Treatment
Modes*
Direct combustion
of char
Indirect combustion
of char
Direct combustion
of char
Indirect combustion
of char
Indirect combustion
of coal
Direct combustion
of coal
Hydrogen generation
by gas reforming
Hydrogen generation
by char oxidation
Hydrogen generation
by gas reforming
For hydrogen
generation
For synthesis gas
generation
Case I
Case 2
Case 3
Case 4
Combined cycle case
Fuel uas case
Sulfur Content ('Pi
In
Untreated
Feed*
1.60
1.60
190
:.9o
0.59
0.69
0.69
0.51
3.93
0.50
0.50
3.30
3.14
6.20
6.:o
1.70
1.70
1.70
1.70
0.60
0.69
4.14
4.14
4.87
1.08
0.67
In
Product
Fuel
0.00
0.00
0.00
0.00
0.00
0.00
—
0.00
0.00
o.o:
o.o:
o.::
1.17
0.00
0.00
1.66
0.98
0.48
o.:o
<0.01
O.I
0.0:
0.02
0.08
0.05
0.04
Disposition of Sulfur in Feed ('')
As
Elemental
Sulfur
74.4
85.6
91.3
96.9
77.1
95.9
-
95.1
98.1
:9.4
40.6
84.:
77.3
98.0
98.4
12.4
51.6
7-5.7
89.:
19.1
91.95
95.7
94.9
97.81
9:.:7
90.9:
To
Atmosphere
13.6
14
7.3
1.7
16.5
1.9
-
:.(.
0.8
o.:
r.o
0.8
0.8
:.o
1.6
O.I
0.5
0.8
0.9
1.1
0.90
o.:o
15 (60)
:4 (95)
23 (91)
C7)
15 (60)
33 fl3ll
16 (631
:4 (95)
14 (56)
13 (5:i
Utility Financing Method.
c/106 Btu (C/106 kcal)
15 (601
19 (75)
17 (67)
:: (87)
5 CD
16 163)
i: (481
14 (561
14 (54)
4 (14)
4 (15)
14 (561
3 (11)
: i 6i
: i 9i
-
-
—
—
5 (19)
10 (40)
23 (91)
i: (48i
19 (75)
II (44)
10 (40)
* Percent sulfur calculated on an as received basis.
* See appropriate process synopsis.
++Based on product heat content.
-------
The results presented in Table 1-2 must be interpreted with care. In
particular, the reported pollution control costs cannot be divorced
from the feedstock indicated; therefore, it is not valid to compare
costs from one process to another unless the feed is the same. Even
with the same fuel, the comparison is of doubtful value. These future
clean-fuel plants will be much'like present oil refineries and no two
oil refineries, even though operating at similar production rates, are
identical. The differences between the basic processes considered in
Table 1-2 are even more pronounced.
There are several other cautionary notes concerning the results
presented in Table 1-2 which are worth emphasizing:
All of the results presented are based on analyses of
hypothetical plants in which the feedstock, unit operations,
and end-product mix were specified in an arbitrary
fashion. Hence,the results cannot be taken to reflect
optimal conditions, either from a. sulfur recovery or cost
viewpoint.
Not all of the processes selected for consideration are
equally efficient or cost-effective in removing sulfur from
the product or by-product streams generated. A variety
of processes were intentionally chosen to illustrate
different alternatives. Hence, the results in Table 1-2 may
reflect more the choice of control units than the intrinsic
emission levels of the basic clean-fuel process.
In general, pollution control costs will be higher when the
feedstock impurity level is greater. Hence, the emissions
levels and the cost estimates in Table 1-2 should be
correlated with the sulfur content of the feed.
It is reasonable to expect that as individual plants are installed and
control processes are carefully engineered to deal with the pollutant
concentrations actually present in the waste streams, both net emis-
sions levels and costs would decrease. Hence, the results in Table I-
might well reflect pessimistic estimates of pollution control costs.
These estimated costs are typically 15 to 30 percent of the projected
cost of manufacturing the clean fuel.
The primary emphasis in this research has been placed on con-
trol of emissions to the environment from the product and waste
1-10
-------
streams indicated earlier in Table 1-1. Technological solutions are
available or are being developed to deal effectively with these prob-
lems. But there are other environmental impacts associated with a
production scale facility (for example, 7.1 x 10°m^/day, or 250 x
I06ft3/day, of pipeline quality gas) that are not readily amenable to
technological fixes:
Community Impact; The total labor requirement for the
plant and mine may range between 1600 and 2400 persons;
related employment might bring the total employment
level to 4000, which implies a community of 15, 000
people. Not only will this influx of people to a new plant
site have a major impact on the local economy, but it
will also have an impact on the environment.
Aesthetics: The physical magnitude of clean-fuel plants,
particularly those gasifying coal, is significant (several
times larger than a major chemical plant producing
ammonia, for example). The chemical processing trains
(reactors, pumps and piping) might be over a thousand
feet long. The boiler plant required to produce energy
for the process would be equivalent to a 300 megawatt (MW)
power station. Furthermore, the mine required to feed a
major coal gasification (or shale oil) facility must be very
large. It is estimated, for example, that the mine at the
El Paso Four Corners site will be one of the largest in
the world. The coal storage pile at any such facility may
have to be as large as 9 meters (30 feet) high, 90 meters
(100 yards) wide and 1600 meters (1 mile) long to provide
a sixty-day supply. The entire plant site might require as
much as 500 to 1000 acres and include large ponds for
cooling water and waste disposal.
Even if the maximum requirements are not met, the potential impact
of clean-fuel processes in the future can be significant. According to
the Final Report of the Supply-Technical Advisory Task Force,
Synthetic Gas-Coal, April 1973 for the Federal Power Commission,
the total projected number of coal gasification plants alone in 1990,
will be 36 based upon projected commercial investment. If, however,
the goals of Project Independence are to be met, it has been estimated
that the 41 clean-fuel facilities listed in Table 1-3 must be constructed
immediately. In order to compensate for decreasing production in
domestic gas fields and to reduce the requirements for imported oil
1-11
-------
Table 1-3
Synthetic Fuel Plants Recommended
for Project Independence
Number of Plants
Product
Quantity
5
12
16
total 41
Shale oil
High-Btu gas from coal
Low-Btu gas from coal as fuel
for power generation
Motor fuel and clean distillate
fuel oils from coal
De-ashed coal or syn^crude
from coal
Fuel grade methyl alcohol
from coal
16 x 106 liters/day (100,000 bbl/day)
7.1 x 106 m3/day (250 x 106 ft3/day)
800-1000 megawatts (MW) of electricity
16 x 106 liters/day (100,000 bbl/day)
16 x 106 liters/day (100,000 bbl/day)
18,000 m tons/day (about 20,000 short tons/day)
to acceptable levels, as many as 165 major synthetic fuels plants
would have to be constructed before 1985. The net environmental
burden of these plants would be substantial even if plant emissions
were controlled as adequately as appears possible from the results
' of this study. Alternative uses of raw fuel would likely result in net
emission levels and environmental impacts far exceeding those of the
clean-fuels plants.
3. ORGANIZATION OF THE REPORT
This report has been divided into two sections as follows:
Section 1: In addition to this introductory chapter, this
section contains chapters
Presenting the approach used in conducting the study
Dealing with three topics common to all process
descriptions
Characterization of pollutants emitted
Nature of control process alternatives
Cost accounting procedures
1-12
-------
Section 2: This section presents the process synopses
for the 12 processes selected for detailed study. Each
process description is presented according to a uniform
format and each can be extracted from the report and used
as an independent guide to the particular process and its
pollution control problems.
Preparation of the report has been guided by two requirements:
The need to have specific information on pollution control
alternatives and costs, particularly for sulfur control, to
aid in the development of reasonable source performance
standards for clean-fuels plants
The need to have a standard reference for use in evaluat-
ing Environmental Impact Statements for clean-fuels plants.
If it is recognized that much of the information presented herein is
preliminary and subject to modification and that many basic assump-
tions, estimates and judgments have been required to compensate for
lack of empirical data, then the report can serve both purposes ade-
quately.
1-13
-------
II. SUMMARY OF ANALYSIS METHODOLOGY
There is a large number of clean-fuels processes which are
currently available commercially or are under development and will
be available for production use in a few years; the number of alterna-
tive pollution control systems which could be applied to these basic
processes are equally large. To conduct detailed engineering analyses
of all possible combinations of processes and clean-up units and to
determine the optimum systems from both an economic and pollution
control viewpoint would have been well beyond the scope of this study.
It was, therefore, essential to adopt an approach which would both
meet the objectives of the program and permit a reasonable degree of
technical detail to be provided for each of the clean-fuel processes
considered.
The approach which was selected consisted of three basic steps:
Classification: All the clean-fuel processes which are now
in use or which could be commercially viable within ten years
were identified and classified. The classification scheme
adopted permitted the various processes in each feedstock-
product class to be compared and contrasted.
Selection; Criteria were established for the selection of a
few representative processes from each feedstock-product
class for detailed study. A total of 12 processes were
selected based on these criteria.
Case Study Analysis: Each of the processes selected was
analyzed in detail as a case study. Alternative pollution
control processes were selected for application to the
basic clean-fuels processes when specific recommenda-
tions were not made by the process developer. This ap-
proach permitted a variety of different systems to be con-
sidered in reasonable depth.
The purpose of this chapter is to present the specific results of
the classification and selection steps and to summarize the method-
ology used in the case study analyses. The discussion is divided into
four parts:
II-1
-------
A review of the classification scheme
A summary of the criteria used to select representative
processes and the rationale for each selection
The approach used to identify and assess pollution control
problems associated with each of the processes selected
A summary of the approach used in preparing the synopsis
of each selected process, including flow sheet preparation,
energy and material balance calculations, description arid
analysis of pollution control procedures and calculation of
control costs.
1. CLASSIFICATION OF CLEAN-FUELS PROCESSES
The first step in the study was to identify all of the clean-fuel
processes which have been discussed in the literature and classify
them in a systematic way. Since only those processes which can be
expected to be commercially viable within the next ten years were of
interest, some classes of processes were eliminated from considera-
tion, including:
Processes which are old and outdated due to recent tech-
nical advances (e.g., Sun-Thermal and Koppers-Kontalyst)
Processes only recently announced on which preliminary
research and development work is now being undertaken
(e. g., EXXON process)
Processes which, although licensed, have major develop-
ment problems to be solved and hence may not be com-
mercially available within ten years (e. g., In-situ thermal
process for tar sands and the In-situ process for shale oil
extraction).
In addition, processes essentially identical to established ones, though
carrying different names, were eliminated. This initial screening,
however, left nearly 50 distinct processes to be classified.
The intent of the classification scheme was to establish the
similarities and differences between processes and to develop a basis
II-2
-------
for selecting representative processes for detailed analysis. Hence,
the processes were first grouped into eight main categories according
to the feedstock-to-clean-fuel-product conversion being accomplished:
Coal to gas
Coal to coal or oil
Lightweight oil to gas
Middleweight oil to gas
Heavyweight oil (and tar) to gas
Oil to oil
Oil-bearing minerals (tar sands and shale) to oil
Oil-bearing minerals to gas.
Then, for each of these categories, a series of generic process steps
(typically unit processes) was established; these served as headings
in the classification matrix. The processes identified in each major
category were described synoptically in the classification matrix for
that category. In addition to process step descriptions, the matrix
includes comments on the status of development and on principal
pollution control problems.
Tables II-1 through II-8 present the process classifications de-
veloped. They provide a review in summary form of all the processes
which have near-term commercial potential and permit a straight-
forward comparison of the features of the various processes. The
processes are not arranged in any particular order in the matrices
nor are the descriptions intended to be comprehensive; however, the
matrices do provide a convenient reference guide to clean-fuel con-
version alternatives.
2. CRITERIA FOR SELECTION OF PROCESSES FOR DETAILED
ANALYSIS
From the conversion processes included in the classification
matrix, 12 were selected for detailed study and analysis of pollution
control procedures and costs. The primary criteria used to make
these selections were that:
The processes, taken together, represent all the major
elements of fossil fuel conversion technology now being
practiced
-------
Table n-1
Coal to Gas Conversion
^5%
Kopp«r»-
(Koppers
Co. '
schift ,'ue r
Chemitechr.ik
mb H)
Research)
CO, Acceptor
Coal Co. 1
and
aiih air and
bed at 315°-
400°C.
C.-.r. -J?e »U do-
Fi«,«.l.i,«
me^ic'cMlf
It requlrtd.
Limited to more
coal).
(.asificalion
and
r.^,^,.,..^,.
Iligss portion of lh«'
procrjf C; round, dried
coal i] (lurried with light
fluidiiei bed hydrogasifier
operating ai 70-105 kg'
gssifier." wygisifir-r or 1
ch.r from :r.o iect.-'i i-.scl
of the Hygss --::. products
hydrogen-rich gi» which i*
supplied for gasification.
U 31JOC.
Mixture of coal and (>2 if
Operation i, a! ,rnlo»p::er-
icpreifjr... KM' CM
temperature is 14BO°C.
fn-r. A rt-voli-ine arate
rrir.oi-inc ash. Operating
^'.""m^d'ssjV.
,.«. Tt, rcnltln, CM,
rer-oving a«h. Operating
pressure it TO kg/cm2.
92T«C.
In a devolauior tection. coal.
dolomite react to itipply beat
ICaO - CO2> to luppon the
reactions; tbe dolomiie is
regenerated by combusting spom
char: hydrogen- rich gag is pro-
duced in the gasifter by reaction
supplier heal required.
Qurncr.inp
ard »..(«r»-e
H^.rrd
cJndonl'a;..
regulaird
portion for
subsequent
tion.
Wilh water.
Acid-t-as
";CoC,
ii-cas pror,,s
rrquir.-s ai-id-
b-j\ required
na.ion.
prior to CO-
but required
Required prior
-^r
,V „„*,,,
Collowed bv an
Hcqjirt-d
Required
rO-Shifl
To'".;,:^,
obtain ar. 11 •
CO ra'id of
Hcqutrt-j
shift rrac;ion .
Not required
for N -diluted
gas
lOO' of erurie
B»i il »ub-
CO-thift reac-
NCK rcqutrrd.
Traci- Salfur
nation catalylt.
Require
Required
I™.«M,
Methanation
All the methane is
gasifler crude ga*
contains CO- H,
More methanatton
required for low BTV
"•"""•
Required.
Dehydration
Extensive drying is
CO2 i* also removed.
Required.
*
Required.
Technology
In pilot plant production
(6S m tons of coal -
42.000 n-.'/dayl
In commerc.al productton
(400-450) BTl' tat (vithout
methanation overieai.
1108 m ton /day -
Pilot pUnt startup in process.
(3G m ton kcal/day -
ST.OOOm'/day).
c«.™-.. „«,„,„
For the pretreatmenl of e«king coals, sulfur existing in the pr«-
ireatrrent off-gas must be withdrawn. As with most processes.
the amount of sulfur remaining in the spent char i* dependent on
characteristics of original feedstock «nd the condition* of hydrogen
Very high off-gas temperature precludes !he formation of any com-
the ash. High off-ga* temperatures should reduce tars, amines.
Sulfur treatment of the regenerator off-gas is required.
-------
Table II-l
(Continued)
^5:-....
• -1.- i'
.1. .-...,. p^. , ...
7Hr
. ...--. \.. ,.„.„,..-
' .'''."'C'A.':?.'-."'..
,...-.. ... ... ...
•-'
,„:-.-.«,,.. -r-
;:..'.;.•.;;••• "*"
,.I3C,
T op
0
pr-fH -0 -P
cm
0
^
•«. ed o
fl OIO°( an
P'
e 0
o K( d a
,..,_ (._
!fl n-a-cr
si rubl.er
H--q4:i-et:
tt'"i-
,*.I«T. II,-
r^-1-..^'i ar
.rfl ,(T ..4
f£H!
Required
.And-Ca*
Sot rtqu:r*dp.'iOr!0
rrquircd bcror* mrth-
Ilt-qu.rril
11. "T«i
wcifcanaiiwiml-.-.
rn::a::u .n th* sail
n:f..-3nau.>R oniv
pt ra.jrt- pro, f SB
a!,o ,,,r.d,rd
pro,,,,,, ou, 9.
„_
Required
•SE,
Kcquirrd
Ftrq-jirrd
and removed t>v
pro.-«S
In novrl and-ws
i leanup. elemental
Oth.nrm <•!.„.
u"d
Rrq-Jircd
CO-3iifl
671-- of the pa*
b--pa»»*s thii
shin reaction
Not required
Required
Trace Biitur
RKju.rfd
Required
Not requ-rcd
Not req^red
Required
„-—
Required
Required
-f.ensivr l!an
\m req-nred
\ot i-fqajrfd
Required
IMivdmiMn
Same as Hv^as
Required
req-Jircd ihan for
Wet. ho-, ion Htu
'-•"' 's -c™ :, ,ntmd«-«! ,.•-« It. «,,,r:,r. .ndurn
-------
Table II-2
Conversion of Coal to
Coal or Oil
^\^Process
^^^Jiteps
P rocesse^\^^
Char Oil Knergy
Development
Process-COED
(FMC)
Synthetic Oil
Process
(I'. S. Bureau
of Mine.-)
Solvent Refined
Coal Process —
(SRC) (The
Pittsburgh &
Midway Coal
Mining Co. )
H-Coal
(Hydrocarbon
Research, Inc. )
Pretreutment
Coal is crushed, dried,
anu dust is cycloned
out
Coal is crushed, dried.
and slurried with re-
cycled oil.
Coal is pulverized,
dried, & slurried
with hot solvent
(290°-430°C)
Coal is dried (at
65°-93°C in a N'2
atmosphere), pul-
verized, and slurried
with coal-derived oil.
Dissolution or Pyrolysis
Coal is pyrolyzed in four stages.
The temperatures of the stages
are 315°, 454°, 538°, and
871°C, respectively. The gas
from the fourth stage flows
counter current to the solids
through tlie other stages. Most
of the volatile products are
thereby collected.
The slurry, mixed with recycled
HQ. flows through the catalytic
fbccd-bed reactor. The reactor
is operated at 450°C and 142 to
250 kg/cm2. The combined
effect of the H., turbulence and
catalyst is to liquefy and
desulfurize the coal.
The coal-solvent slurry is
pumped, together with hydro-
gen, through a prehcater to
a reaction zone or dissolver.
The dissolver is operated at
a temperature of about 430°C
at 71 kg/cm2. About 95% of
the moisture and ash-free coal
is dissolved.
The coal-oil slurry is charged
continuously with hydrogen to a
reactor containing a bed of ebul-
latcd catalyst. The coal is cata-
lytically hydrogenated and con-
verted into liquid and gaseous
products at 160-190 kg/cm2
and 455°C.
Ash & Sulfur Removal
COED process generates a
hot char which can be
further desulfurized by
lime-treating or gasific-
ation (COGAS)
The separated oil is
centrifuged to remove
ash and inorganic
sulfur.
After separation of excess
H2, the dissolver effluent
is filtered and ash and un-
dissolvcd coal is removed.
Sulfur can be recovered
from the stack gas by a
• modified Claus process.
Ash is removed in the
fluid coking section.
HQ and KH-j arc removed
in the form of sulfuric
acid and ammonium sui-
fate.
Removal of Solvent
and/or S eparatign
of Product Oil
The volatile products are con-
densed and the product oil is
hydrodesulfurized prior to
sale. Noncondensable fuel
gases are desulfurized and
consumed on site.
The product from the reactor is
cooled; oil and gas is separated.
The separated gas in the high-
pressure receiver is recycled
after HjS and NHj is removed
by absorption in water. Part
of the separated oil is recycled
to convey more coal through the
feed pump into the reactor
Filtered oil is heated and vac-
uum-flashed. The temperature
and pressure of the vacuum
drum is controlled to provide a
bottoms product of 93°C to
187°C melting point. The over-
head product is further frac-
tionated into aromatic products
and a wash solvent. The bottom
fraction is a hot liquid, which is
the major product of the process
The reactor effluent is separated
into gas and liquid streams. After
removal of light HCs from the gas,
it is recycled. Liquids contain-
ing the coal residue are sent to a
fractionator. The heaviest liquid
containing the coal residue is
steam-stripped to remove light
HCs, and then vacuum-distilled.
The vacuum bottoms, which are
reduced to 50% solids, are
pumped to fluid coking.
Technology
A pilot plant (33m
ton/ day) is in oper-
ation. In all coal-
to-oil conversions
illustrated, syn-
thetic crude can be
converted into
product oil by con-
ventional refinery
processes
A 1/2 m ton/day
pilot plant is in
operation and
another pilot
plant is planned.
A 45 m ton/day
pilot plant is
under construc-
tion. Sol vent -
refined coal can
be hydrotreated
to product liquid
hydrocarbon.
Beingtestedon
process devel-
opment unit
(2-5 m ton/day).
Comments on
Pollution
Sulfur requires
removal from
product liquid
and gas streams
and from char
Product oil must
be hydrodesul-
furized. Char
contains sulfur.
For coal products.
some sulfur re-
mains in coal;
for syn-crude
products, sulfur
canbe removed •
by hydro desul-
furization.
Residues contain
sulfur andpyrites.
Problems are
similar to those
of the Synthetic
Oil Process.
-------
Table II-2
(Continued)
^^^^ Process
^\_ Steps
Processes v^
^^
Consol Synthetic
Fuel Process-CSF
(Consolidation
Coal Co. )
Methanol Synthesis
Fischer-Tropscb
Process
Meyers Process
-TRW
Ledgemont
Process
-Peabody Coal
Pretreatment
Coal is crushed, dried,
and slurried with a
solvent. Large coal
particles are dis-
charged.
Coal is oxygen gasified
to synthesis gas.
Coal is oxygen gasified
to synthesis gas.
None
None
Dissolution or Pyrolysis
Slurry is pumped through tubular
furnaces where it is heated to
extraction temperature. Extrac-
tion occurs, however, mainly in
an extraction vessel.
Indirect conversion of coal to oil
through synthesis gas. Not a
dissolution or pyrolysis process.
Indirect conversion of coal to oil
through synthesis gas. Not a
dissolution or pyrolysis process.
Selective dissolution of pyritic
sulfur by ferric ion.
Selective attack on pyritic sulfur
by oxygen under pressure.
Ash & Sulfur Removal
Untreated coal is removed
in hydroclones. H_S from
gas treatment is converted
to sulfur in a modified
Claus plant.
Similar to Texaco
process.
Similar to Texaco
process.
Under development.
Under development.
Removal of Solvent
and/or Separation
of Product Oil
By fractionating the product
from the reactor
Catalytic conversion to
methanol and condensation.
Catalytic formation of oil
from synthesis gas.
I nder development.
Under development.
Technology
Pilot plant
64 m ton/day
was operational
for months.
Commercial
overseas.
Commercial
overseas.
Bench scale
under develop-
ment.
Bench scale
under develop-
ment.
Comments on
Pollution
Problems are
similar to those
for the SRC
Process.
Similar to
coal gasifica-
tion.
Similar to
coal gasifica-
tion.
Not yet
analyzed.
Not yet
analyzed.
-------
Arable II-3
Conversion of Lightweight
Oil to Gas
Process
^s. Steps
^X^
Processes^.
Catalytic
Rich Gas -
CRG Process
(British Gas
Council)
JGC Methane
Rich Gas
Process
(Japan Gas-
oline Co. )
Gasynthan
Process
(Basf/Lurgi)
Sun/Thermal
Process
(Howe-Baker)
ADTEK
(IGT)
Prefraction-
ation
Feedstock is
distilled to re-
move all heavy
factions (BP
over 204°C
These will de-
activate HDS
catalysts.
Required.
Fractions up
to 260°C
allowed.
Allowable may.
to 215°C BP
Uses
naphtha
Takes light
hydrocarbons
Hydrodesul-
furization
Mixture of oil and
H2 reacts in pres-
ence of catalyst
and converts sul-
fur compounds to
H2S. Numerous
processes avail-
able (e. g. .Hydro-
fining). For higher
sulfur contents in
fuel, convention-
al HDS units are
ised.
Required
Required
Required
Molecular sieve
treatment and
hydrodesulfuri-
zation is re-
quired.
HgS
Removal
In a bed of
ZnO. ZnS
is regen-
erated and
sulfur re-
moved.
Required
Required
Required
Required
Vaporization
Mixed with steam
and superheated.
This prevents car-
bon formation and
supplies heat for
reforming reac-
tion.
Required
Required
Required
Required
Low Temperature Catalytic
Reforming
Oil and steam react in pres-
ence of catalyst (at 550 psig)
to produce crude gas. 7ne
steam is the source of H~
required to assure correct
H2/CO ratio. Exit gas tem-
perature is 400°-480°C.Then
gas is hydrogasified (to
increase CH4 content). Some
crude gas is used to manu-
facture H2 for HDS and oil
hydrogenation.
Oil and steam are continuous-
ly introduced. Continuous
reaction permitted by a
unique catalyst.
Reaction occurs in adiabatic
reactor at 400°C. The crude
gas is fed to a 2nd adiabatic
reactor where H2+CO2 react
(over a catalyst) to increase
CH4 content.
Vaporized naphtha contacts
flame and undergoes imme-
diate endothermic thermal
decomposition. Air is used
instead of O2-
The mixture of steam and
hydrocarbon feed (at 438°C
and 26 kg/cm2) is fed to the
fixed bed (510°C) catalytic
reformer.
Metha-
nation
Required
Required
Required
Not re-
quired
Required
Dehy-
dration
Required
Required
Required
Not re-
quired
Required
C02
Removal
Required
Required
Required
Required
Required
Technology
Low BTU plants
in existence.
Hi BTU plants
will begin opera
tion in 1973.
Naphtha is in
short world
S'.ipplv.
In commercial
use.
In commercial
use to produce
low BTU gas.
Hi BTU produc-
tion in pilot
plant stage.
In commercial
use.
Tested on expe-
rimental scale.
Comments on
Pollution
There is little sul-
fur disposal problem
due to low sulfur
content of feed-
stock.
There is little sul-
fur disposal problem
due to low sulfur
content of feed-
stock.
There is little sul-
fur disposal problem
due to low sulfur
content of feed-
stock.
There is little sul-
fur disposal problem
due to low sulfur
content of feed-
stock.
There is little sul-
fur disposal problem
due to low sulfur
content of feed-
stock.
oo
-------
Table II-4
Conversion of Middleweight
Oil to Gas
^^^^ Steps
^\_^
Processes ^^^
^"^v.
Chemically Active
Flu id Bed Gasifica-
tion -
CAFB (ESSO)
Gas Recycle Hydro-
genation — GRH
(British Gas Corp.
NRG Reforming
(MiFuji Iron
Works Co. )
Petrogas Process
(Gas Machinery
Company)
Oil Hydrogasi-
fication
(IGT)
Pretreatment
Required
Feedstock + H2.
Max BP is 345°C
Heated by re-
cycled gas.
Any hydrocarbon
fraction with an
end boiling point
of 370°C or less
(such as naphtha.
kerosene). Air
and hydrocarbons
are preheated.
Feedstocks
heavier than
naphtha.
Straight -run
distillate feed-
stocks with
boiling point up to
360°C. Feed-
stock and Hn
are preheated.
Gasifier
Oil is gasified with air
at 815°C in a fluid bed
of limestone. All sul-
fur can be removed as
CaS . The limestone is
regenerated & recycled
to gasifier.
Hydrogenation occurs
at about 760°C and
71 kg/cm2
Hydrocarbon, air and
steam are fed to the re-
actor, which operates
at about 815°C. Some
of the product gas is
recycled.
Feed undergoes ther-'
mal reforming with
steam.
Noncatalytic reac-
tion of oil with
free hydrogen at
elevated temp.
595°-760°C and
pressure 36-107
kg/cm2.
Aromatics
Removal
Required
In Benzol absor-
ber.
Not required.
By oil.
Aromatics are
removed in
Benzol absorber.
H.,S Remova
Required
Required
Required
Removed
by Stret-
for Process.
Required
Secondary
Hydrogenation
Required
A side stream of gas
is diverted for pro-
duction of H2 added
during preheat.
Remaining gas is
hydrogenated.
Not required
Can be by hydro-
genation or metha-
nation.
Methanation is
required.
Dehydrating
lequtred
Required
Required
May be re-
quired.
Required
Technology
Pilot plant design
stage.
Pilot plant stage
for hi BTU gas.
Low BTU gas
plants are in com-
mercial existence.
Two small com-
mercial plants
are in operation.
In development
stage.
A large-pilot plant
with a capacity of
11,300 m3 /day of
product gas was
operated at IGT
in early 1950's.
Additional pilot
plant work on a
combination of
hydrocracking and
pressure hydro-
gasification was
carried out from
1955 to '959.
Comments on Pollution
Regenerator purge gas stream
contains SO2, vanadian, lime.
race metals.
Additional sulfur in feedstock
can be handled by a con-
ventional acid-gas clean-up
process.
-------
Table II-5
Conversion of Heavyweight Oil
(and Tar) to Gas
>i Steps
proceH.icV
ONIA-GKC1
(Office of
Naltonal
IVl'A/.otc)
Sln-11 Parlial
Oxidation
Tex:«-o I'ur-
ti»l Oxidn-
H-Ons
(Hydrocarbon
K.-search Im-)
HOC Unit
Cnpyn.lysis
Ri-Midimm
(KMC Corp.)
SKilAS
Unier-
Matenuls)
rijri.- I'etn.-
•M-M In.-. )
I'-lui.l Med
C^ilinn - Kllll
(Hrilish C.is
Prctreatmrn
and
All oils
No limit on
feedstocks
Also com 'Ms
hitfh sulfur
res id oil,
SNC
No limit on
fee, [stocks
in IVcdslocks
Tor ln-avy
hydrocarbon
ifiwv liquid
furls.
,'m! or I,.'- ivy
I ses jiny p---
cru.le.
<;asific;ilion
Oil and jtlc:ini in atoni-
cyclinKrcfonmnjiproc-
j;as Irmp'-raU].-.' i.s
870*V. (ippration i« ul
uniliicnt pt'i-ssiirc. f;it;i-
Pr'i-ln-iili-cl r<-i-dsto.-k und
ox.yi!i-n JUT «-h;u-(,'r.l In
tin- non-i'iit.-ilytii- n-
actor opiTUlAd ill 1005°C
iri10°{' iinii up lo 105 kjj/
cm2
Fi.(.(HKick is !iyilrti«;is.
iTii-d HI tin- prrsrn.'r nf
H2 i.. piiKlur.- lluht hy-
it (piMi-rt. f"kr
mtd ti.-avy I'-sidn:.! ml
II;, :itnu>sph'-i>- Inr thi-
liyilrnj{:isirii-:ili<>n «lr|>.
Tb.- .-riiili- is ft-ii.'lii.ti-
:it.-cl. Tlic ch-;i,spliiilii-tt
oil and thr vircin 1HS
lo 537°r fractions ;.n:
hydrocracki'd lo .yield
18.r)°C and liRhU-r ma-
Irrhil. This i-:in h.-
used In ni:ihi> KN(i by
thi- CRG prtM-.-ss. Tin-
will. oxy«.-n. Sl.-:im i«
aiiih'd to Khil'i the ^us
to make nioi-c Il2 which
ran !»• ,,s.-d in hydro-
d-a^king.
Prclifiilfd f«'fd nil is fed
lo tin- fluid hi-d hydro-
Kasifi.-r ti> prodiit-i- Cft.,-
rich gas. f)i>pd t-:ir-
lion i« uasifi'-d in a Uasi-
Uii-HiK K«S whi.-h pr.,-
viil<-K !lz U. tin- hydn»-
Casil'iiT.
(427 - 4B2°C) tn remove
volatility :.nd pass.'.! on
lo (In- I'nu-tii m:i IOT-. T.-mp.
in thi- :ird ami 4lh slaves
:n-i> r,:tn - 6r>n(1r and
870r'C r<'np. Casc.s
lyclt-d luirktlirinichlhi'SiM-
ions hy(ifO['a''lion >'onipo-
in-uls uf Ihi- f'-i-dslr>,-k
become .sui-ci -ssively
rrachi'd as UK- niai>TJal
ndvan. <-K tlii'»u^li I1i<- e:<-
ii-rtiiilly hi-aii'ii hi'H. Coke
ih:il forms on the ln'd r<-
acls with sU-a.,1. The
|.rt\
Tied :il this sl:if;e. Coki<
lirsl since go lo another
ill 1 100 - 1200" C. . pro-
.Itii-int! additional Fiii-l jpis.
l'rehe;>le,t Teed is alnm-
.,F i-nkr particles ;.nd
flni.li?cd l>.v ii M2-ri.:h
ii op.Tiih'd at 7GO"( '
;,.nl (i-I kR/cin2.
Oucrifhiiii!
Wiilct- wash
cooler s.-r-iitil>.-|-.
qu«-»M- «Ul. :.
iydi'ocasifii-1-
walei-Scnililiers.
rMed." ^ P;'
w.ishi il IA i li
Is.'S ini.'i'
.\rorn:ilti:s are
hetwl i»nrlicr
Al "' Siilhit CO shill 'l^ie- HyS M.-tl,;,. i,,.hy-
•1 "!1 ii >'1" '"Vl'rv Ilt':" ''"" U*-!""v:'il "Jitit.ii dtMtion
Sam.- as CKC process.
Same as c.ial uiiMifirtitinn.
Same as coal Kasilic:.lion
Same as coal (."sitValion
Sunie a» <-oal i:asifir:n.
S:irne as for co;il c:>siric:.(for)
l).'v..|(,|M,i(: slaj.e.
l.i mn r i- 1 i.-
Mas passed ilspMol '
l>lani st!i(>r and
ready I'or demonMT-a-
i>l»iii;
In developini; sta^e.
I'ilflt ptilllt Sla(.e.
I.:irtfi- pilol pU.nl
No large plants have
coniniriciul unit with
ti 140,000 in /day
capacity has hccn
opri-atfd In .l.npan.
Can hr used for heavy
nil ensificatlon.
<• '
.Sulfur removal prolilt-ms ore
uasificalion.
similar to those for coal
UasUication.
Sulfur removal problems an-
Simitar to those for coot
Kasiricullon.
Sulfur removal problem H are
similar to ihoRc for coal
Casiftcution.
Sulfur iL-moviil problems
arc similar tn thoue for
coal fill Bin cation.
HIT similar to those for
roal eutiifiriition.
Sulfur ri-moval probli;mn
are simiLir to those foe
coal cuRlfioalion.
11-10
-------
Table II-6
Oil to Oil Conversion
"^Process Steps
^^.
Processes \~
Hydrofining
(ESSO)
RCD ISO-MAX
(LOP) and
RDS ISOMAX
(Chevron)
Residfining
(ESSO & Union)
H-Oil Process
(Cities Service
and Hydro
Research)
HDS Process
(Gulf and
' Houdry)
Feedstock
Limitations
Can use resid
(distilate) oils.
Mix Oil and H2
Mixed with H2Tich
gas bled from reactor
and heated.
Required
Preheated in furnace
and mixed with hot
H2-rich gas.
Required
Required
Hydrodesulfurize (P+T)
In fixed bedjregenerable
catalyst. The liquid
product stream is
passed to separators
260-425°C and
5-57 kg/cm2
Catalytic desulfurization
in fixed bed.
71 kg/cm2. Oil is recov-
ered in high and low, hot
and cold separators.
.Most sulfur is removed
by a hydrodesulfurizer.
With cobalt-moly in
ebbulent bed. Demetala-
tion and desulfurization
reactors are in series.
Operates at 212 kg/cm2
Fixed bed reactor.
57-212 kg/cm2
HoS Removal
In a stripper
column.
Acid gas clean-
up of circula-
ting H, re-
quiredt
Required.
Required.
Required.
Required
Sulfur
Recovery
Required.
Required.
Required.
Required.
Required
Technology
In commercial use.
For all oil to syn-
crude conversion
processes, sulfur
removal becomes
increasingly diffi-
cult for heavier
factions due to
higher metal content.
In commercial use.
Reactor plugging is
necessitating addi-
tion of a quench
system.
In wide commercial
use.
In commercial use. .
In commercial use.
Comments on Pollution
For low sulfur oils (lighter factions), H2S is
removed by ZnO scrubbing. For high sulfur
oils (heavier factions). H2S is removed by
amino- scrubbing followed by Claus or similar
process.
-------
Table II-7
Conversion of Oil-Bearing Minerals
(Tar-Sands, Oil Shale) to
Syn-Crude
Process
Steps
Processes
Feed Separation
Retorting
Oil
Technology
Comments on Pollution
Petrosix
(Cameron & Jones)
Finely crushed oil shale
Feed and recycled hot gas
react by pyrolysis to form
oil vapor. Spent ore is re-
moved. The oil vapor is
condensed.
The processing of oil to a
clean syn-crude is identi-
cal to the "Oil to Syncrude"
conversion.
In commercial use
Disposal of large amounts
of waste material per bbl
of syncrude may be
problem.
Tosco- II Process
(Oil Shale Corp.)
Crushed and screened
oil shale is heated to
260°C. Heat is supplied
by flue gas.
Rotating retort with 650°C
hot balls (ceramic balls).
The spent shale and balls
are withdrawn, screened,
separated; spent shale is
disposed of; and balls are
reheated. Hydrocarbon
vapors are separated in a
fractionator. H2S and COg
are removed.
The processing of oil to a
clean 8yn-crude is identi-
cal to the "Oil to Syncrude'
conversion.
This is the only
operating U. S. shale
process. Similar
technology is exten-
sively commercialized
by the USSR. The
Tosco Process has
been demonstrated at
a rate of 907m ton/day
Disposal of large amounts
of waste material per bbl
of syncrude may be
problem.
Great Canadian
Oil Sand Co.
Process-G COS
Tar sand and water is
heated in rotating drum.
Gravity separates bi turn in
and sand. Bitumin is
dehydrated in centrifuge.
In 3. delayed coker; hydro-
desulfurized to yield H2S.
The processing of oil to a
clean syn-crude is identi-
cal to the "Oil to Syncrude'
conversion.
In commercial use.
Disposal of large amounts
of waste material per bbl
of syncrude may be
problem.
In Situ Thermal
Process for Tar
Sand
Tar sand is burned, in place, by forcing air into an
injection well or heating with steam. Hydrocarbon vapors
and bitumins are forced and collected in a. production or
recovery well.
The processing of oil to a
clean syn-crude is identi-
cal to the "Oil to Syncrude'
conversion.
None is in commercial
use. The reaction is
hard to control.
Eliminates spent sand
handling problems.
Oil shale is fractured by pressure or explosion. Heat
for retorting is accomplished by partial combustion of
some shale with injected air, or other technique.
In Situ Thermal
Process for Oil
Shale
The processing of oil to a
clean syn-crude is identi-
cal to the "Oil to Syncrude1
conversion.
Conceptual phase
Underground retorting
would avoid air pollution
control.
Gas Combustion
Retort Process
(U. S. Bureau of
Mines)
Crushed shale
Shale falls successively
through preheating, retort -
ing (450°C), combustion
(927°C) and cooling (55°C)
zones in a cylindrical retort.
Air is supplied for combustion
The suspended oil mist passes
through centrifugal separators
and an electrostatic precipitator
The processing of oil to a
clean syn-crude Is identi-
cal to the "Oil to Syncrude"
conversion.
Has been demonstrated
on 330m ton/day scale.
An undesirable trait is
that oil products drip
back into the hotter sec
tions where it ia re-
cracked.
Disposal of large amounts
of waste material per bbl
of syncrude may be
problem.
Union Oil
Process for
Shale Oil
(Union Oil
Co.)
Crushed shale
Shale enters and is pushed up
from the lower end of three
vertical truncated cones
against down flowing process
gas. In a 4th retort, carbon
residue in the shale is com-
busted with air (1205°C).
Its product gas is separated
from the oil and burned to
preheat the N- -rich recycle
gas from the other three
retorts (510°C)
The processing of oil to a
clean syn-crude is identi-
cal to the "Oil to Syncrude"
conversion.
Has improved thermal
balance as compared
to the Gas Combus-
tion Retort Process.
Has been demon-
strated at a rate of
91 m ton/day.
Disposal of large amounts
of waste material per bbl
of syncrude may be
problem.
-------
Table II-8
Conversion of Oil Shale
to Gas
^Process
^^^ Steps
Processes^
"\.
Hydrogasifi-
cation of Oil
Shale (IGT)
Pretreatment
Crushed
Many othe
Hydro gasification
Crushed shale is fed to the top of
a three stage hydrogasifier. In the
1st stage shale is preheated
(538°-705°C) by counter cur-
rent contact with reaction product
gas. Preheated shale goes to a
reaction zone where hydrogasifi-
cation occurs. The spent shale
passes through a 3rd gas preheat
zone where spent shale serves to
preheat the cold feed hydrogen.
The pressure of the operation is
in the range of 19-29 kg/cm2.
Quenching
Product from the hydrogasifier
is quenched and liquid HC and
gas are separated. Liquid
HC can be fed back to the re-
action zone of -the hydrogasi-
fier.
Purificat ion & Methanation
After quenching, the product gases
are scrubbed to remove CO, and
HgS , and can be shifted and
scrubbed again or methanated. The
remaining gas, composed princi-
pally of H? + CH4, can be sepa-
rated into IlQ-rich gas and CH . -
rich' gas by cryogenic separation.
Hg is recycled to 'he hydrogasifier.
The CH4 is of pipeline quality.
Technology
Proven on bench
scale. Pilot plant
is planned.
Comments on Pollution
Pollution problems are
similar to those for coal
gasification. Large
quantities of spent shale
must be disposed of.
roil shale togas schemes are available. They can be derived by first converting feedstock to oil and then
to a gas. Numerous combinations of "oil shale to oil1' and ''oil to gas" conversion processes (discussed
separately) define the state-of-the-art.
-------
Each clean-fuel conversion mode addressed in the classi-
fication matrix be represented (e.g., coal-to-gas, oil-
to-gas)
The process be commercially viable or likely to reach
commercialization within ten years
The selected processes exhibit the major pollution con-
trol problems expected to be faced by the fuel industry.
These should include processes exhibiting novel pollution
problems or control techniques.
Owing to the relative abundance of coal as a feedstock and the rapidly
growing interest in coal gasification conversion processes, a dispro-
portionate number of process selections were made from this category.
Consequently, the planned investigation of oil-to-gas conversion pro-
cesses was limited to the heavier oil feedstocks, omitting light and
medium weight oil conversion schemes, since these heavier oil con-
version processes are of greatest concern relative to removal of
sulfur from the product streams.
The processes selected for detailed study, and the rationale for
each selection are indicated in Table II-9. Each of these 12 processes:
The Synthane process
The CC>2 Acceptor process
The Lurgi process (High-Btu gas)
The Hygas process
The COED process
The SRC process
The Texaco Partial Oxidation process
Desulfurization of Crude Oil
The Gas Combustion Retort process
The Lurgi process (Low-Btu gas)
The Koppers-Totzek process
The U-Gas process
is discussed in Section 2 of this report.
3. IDENTIFICATION AND ANALYSIS OF POLLUTION CONTROL
PROBLEMS
This section summarizes the approach used in the identification
and control of emissions from clean-fuel processes. The principal
steps outlined are:
11-14
-------
Table II-9
Selection of Clean-Fuel Processes
for Detailed Study
Conversion
Mode
Coal ID gas
Coal to gas
Civl to pas
Coal to gas
Coal to (ias
C'oal to (ias
Coal to (las
Coal to
Coal or
Syn-Crude
Coal to
Syn-Cntdc
Oil to
C'lean Oil
Heavyweight
Oil to Gas
Oil Shale to
Gas and Oil-
Bearing
Minerals to
Syn-Crude
Process Selected
Synlhane
Lurgi (High-Btu)
COi Acceptor
Hygas
Lurgi (Low-Btu)
IHias
Koppcrs-Tot/ek ( K-Ti
(K-T)
Solvent Refined
Coal (SRC)
Char-Oil- Knergy-
Developmcnl(COKD)
Hydrodesulfuri/ation
(HDS)
Texaco Partial
Oxidation process
Gas Combustion
Retort process
Reason for Process Selection
• The Synthane process is typical of modern coal gasification processes
• Pollutants generated arc similar to those from most coal gasification processes
• The product gas cleaning, purification and sulfur recover)' steps are typical of most gas processes
• Information about this process is readily available.
• The Lurgi gusifier is presently commercially available for (he manufacture of high-Bin gas from coal
• Detailed process information is available in the published literature
• It exemplifies traditional gasification technology
• It produces significant quantities of tars and oils which can cause wastewaler purification problems.
• Uses novel technique for supplying the cndolhermic heal retjtvral to sustain the gasification reactions
• Special pollutant removal problems are encountered
• The process is now in the advanced pilot plant stage. It may soon be commercially available
* It can have unusual wastewater problems.
• Hygas is currently the most advanced of the modern coal gasification processes
• Most of the product gas is produced, without a catalyst, by direct liydrogasification
• The Lurgi gasilicr is commercially available
• An installation for the production of low-Bin gas exists
* Serves as an example of a moving fixed-bed reactor
• Exemplifies one of the low-Btu gasification systems being developed using modern technology
• Serves as an example of a fluidi/cil bed reactor and the application of a combined power cycle
• The K-T gasilier is commercially available for the production of synthetic gas. It is readily adaptable for the production of low-Btu fuel gas.
• Serves as an example of a high- tempera In re dilute-phase reactor.
• The SRC Process is probably the most developed of the various coal -dissolution processes. An advanced pilot plant is now being constructed
• This process permits manufacture of either low-sulfur coal or syn-crude. Both options will he discussed.
• Has been in pilot plant production
• It is representative of the type of process which converts coal feedstock to syn-crude by pyrolysis.
• The common process for removing sulfur-containing compounds from hydrocarbon liquids is called hydrodcsulfurization. There are many
similar processes available for hydrodcsulfuri/ing distillate oils and residuals. These can be operated at varying degrees of severity to achieve
differing product oil sulfur content. Because of this complicating factor, the overall subject ol nydrodcsulfurization of fuel oils will
he treated as a separate study. The HDS process developed by Ciulf is representative and has information available. It is described as an ex-
ample of hydrodcsulfuri/ation-typc process mentioned in this study.
• The Texaco process, for the conversion of heavy crudes and residuals, has been installed commercially.
• It has similar pollution control problems to other processes in this category.
• A pilot plant of significant size has been operated
• It is typical of retorting processes
• It addresses pollution problems that might be experienced with the retorting of coal.
• This process can be used to recover oil directly from shale or, through inclusion of an oil gasification process, can manufacture gas.
11-15
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Identification of discharge sources
Characterization of emissions
Definition of control processes.
(1) Identification of Discharge Sources
Many of the selected processes have not been in commer-
cial operation. During the process development and pilot plant
stages, the emphasis was on "making the process work, " rather
than on protection of the environment. The small scale of op-
eration usually meant that any discharges encountered were
readily absorbed within the local pollution control system or
assimilated by the surrounding environment without substantive
or lasting harm. However, pollution control problems are now
emerging as part of the overall effort to scale-up these pro-
cesses to achieve commercial operation.
For these reasons, the available flow diagrams and asso-
ciated descriptions of the selected processes were inspected with
care to identify potential waste streams. In many cases, it was
necessary to apply engineering judgment to project a discharge,
where none was indicated by the process developer.
The sources of significant waste streams are identified
within each process synopsis, both in the process flow diagram
and in the process description. A summary listing of these
sources, the nature of the waste streams, and manner of their
disposition are presented in Table 11-10. The sources are listed
in approximate order of process flow. It is evident that the
sources of waste streams are many and varied. In fact, there
is hardly a step in any of the clean-fuel processes that does not
represent a potential source of environmental pollution. Imag-
inative process design is required to channel these pollution
streams to the best advantage of both the process and the en-
vironment.
(2) Characterization of Emissions
Characterization of the amounts and compositions of
emissions proved much more difficult than the identification
11-16
-------
Table 11-10
Nature and Emissions of Major
Sources of Pollutants
Source •
Feedstock storage
Feedstock cleaning
Feedstock crushing and grinding
Water treatment
Steam generation
Stack gas cleanup
Oxygen generation
Gasification
Wastewater treatment
Ash ponds
Shift conversion
Sulfur recovery
Methanation
Final purification/drying
Boilers and cooling towers
Heat exchangers
Washrooms and work areas
Land surface
Waste Streams
teachings and runoff
Rocks and debris
Coal and shale dust
Coal fines
Wastewater
Stack gas
Ash
Stack gas
Sulfite sludge
Stack gas
Wastewater
Tar
Char
Ash
Spent shale
Treated water
Ammonia and phenol
Hydrogen sulfide
Solid sludge
Leachate, Drainage
Acid-gas
Wastewater
Heat
Spent catalyst
Sulfur
Tail gas
Spent catalyst
Wastewater
Heat
Spent catalyst
Vent gas
Wastewater
Spent iron or zinc oxide
Slowdown water
Steam
Cooling water
Sanitary sewage
Refuse
Surface runoff
Principal
Components
Pyrites
Carbon particles
Coal particles
Dissolved salts
See Table III-2
See Table III-4
See Table m-2
Nitrogen
See Table III-3
See Table III-4
See Table 111-4
See Table III-4
Dissolved salts
See Table III-2
See Table III-3
Cobalt, molybdenum
Bauxite
Nickel
Sulfur dioxide
Dissolved salts
Heat
Organic/nitrogen cpds.
Sediment, oil
Disposition
Treatment
Disposal
Containment and recovery
Recovery
Treatment/utilization
Treatment and discharge
Disposal/utilization
Discharge
Treatment and disposal
Discharge
Treatment and reuse
Utilization
Utilization
Disposal/utilization
Disposal
Reuse
Recovery
Sulfur recovery
Treatment and disposal
Treatment and reuse
Sulfur recovery
Treatment and reuse
Recovery
Treatment and recovery
Storage
Discharge or treatment
Disposal
Treatment and reuse
Recovery
Treatment and disposal
Discharge
Discharge, reuse
Disposal
Treatment and reuse
Discharge
Utilization/disposal
Treatment and disposal
Treatment and disposal
Discharge
Process/Unit
Neutralization with lime
Water washing, filtration
Cyclone separators, bag filters, enclosure
Lime softening, zeolite treatment
Wellman Lord, limestone injection, lime scrubbing, catalytic oxidation
Land filling/incorporation in construction materials
Chevron, Phenosolvan, Phosam, oxidation, biological
Conversion to oil, direct gasification, direct combustion,
indirect combustion
Gasification, desulfurization, direct combustion
Land filling/ incorporation in construction materials
Minefiliing
Claus, Stretford
Dewatering, oxidation, landfilling
Hot potassium carbonate, Rectisol, Sulfinol, MEA,
DffA, etc.
Chevron, Phenosolvan, Phosam, oxidation, biological
Oxidation
Incineration, Beavon, SCOT
Chevron, Phenosolvan, Phosam, oxidation, biological
Oxidation
Oxidation, iron and zinc oxide treatment
Waste heat recovery
Biological
Incineration/landfilling
-------
of their sources. In the case of major waste streams, which
have been defined by the process developer, it was possible to
estimate efficiencies of the proposed control processes. In
other cases, estimates had to be more qualitative, and were
based on engineering judgment.
A detailed discussion of these emissions covering the
nature of emissions, their estimated amounts, the mechanism
of formation, the specific control problems, and the threat they
pose to the quality of the environment, is presented in Chapter III.
The accompanying tables illustrate selected compositions of gas-
eous, liquid, and solid waste streams, the sulfur balance of rep-
resentative processes, and the sources and effects of major
pollutants.
(3) Definition of Control Processes
The definition of pollution control processes was provided,
in some cases, by the process developer and, in other cases,
by the study team. The distinction is indicated by a dashed line
on each process flow diagram. For each clean-fuel process, a
discussion of the relative merits of alternative control processes
and a detailed computation of the incremental costs of clean fuels
attributable to pollution control measures are given. A general
treatment of these topics is also presented in Chapter III.
Table 11-10 lists the major control and treatment processes for
waste streams emanating from each of the major potential
sources.
Four factors were considered in selecting alternative
pollution control processes for application in the process synopses:
Ability to meet applicable emission standards
Suitability for the specific waste streams
Demonstrated effectiveness in similar applications
Life-cycle cost.
Since emission standards for clean-fuel processes have not
been promulgated, similar existing standards were used as a
guide. Assessment of the suitability of control processes for
specific waste streams required both an evaluation of their past
performance in related applications and an extrapolation of their
11-18
-------
sensitivity to variations in waste stream composition. Incre-
mental life cycle costs were computed in accordance with both
the discounted cash flow and utility accounting methods, to ac-
commodate anticipated differences in accounting procedures.
Finally, it must be noted that the proposed control pro-
cesses are only conceptual. Firm decisions on their selection,
sizing, and costs must await detailed engineering feasibility
and design studies, which were clearly outside the scope of this
study.
4. PROCEDURES FOR PREPARATION OF PROCESS SYNOPSES
The principal objective of the research program described in
this report was to investigate alternative procedures for control of
emissions from clean-fuel processes by analyzing in detail the 12
representative processes selected from those described in the classi-
fication matrix. This case study approach provides more detailed
technical information than would an attempt to survey all the classi-
fied processes in a superficial way, but it did hot permit every com-
mercially feasible process to be analyzed. The results presented in
Section 2 of the report can, however, be extended to provide some
insight into the pollution control problems of processes closely related
to those described in detail.
The general approach used in considering the process studies was
to assign each process to a research team. One process (Synthane)
was selected as a prototype, and the analysis of this process was com-
pleted first to establish an example for the research teams to follow.
A uniform format was adopted for the presentation of information for
each of the process descriptions, including the flow sheets, material
and energy balances, pollution control alternatives and costs. This
format not only ensures a consistent level of detail in the analysis
but also simplifies the comparison of results from process to process.
The processes' descriptions are divided into five sections as follows:
•Introduction
Process display
Process description
Discussion of pollution control processes
Costs of pollution control.
11-19
-------
The information contained in each section and the approach used in
preparing it are described in the following paragraphs.
(1) Introduction
Each process synopsis begins with a brief introduction in
which the origin, state of development, and near-term plans for
the process are described. In addition, significant features of
the process which distinguish it from other clean-fuel conversion
methods are indicated.
(2) Process Display
This section of each synopsis outlines the basis for the sub-
sequent analysis and presents the assumptions made and the con-
ventions adopted. The flow sheets and stream compositions are
also given as well as an explanation of its layout format and
symbolism.
The flow sheets were prepared detailing the unit processes
used to convert each fossil fuel feed to a desulfurized fuel.
Special emphasis was placed on describing the handling and
disposition of chars, residuals and sulfur as well as on quanti-
fying the sulfur and any saleable by-products derived from the
pollution controls applied to each process. Where possible,
effluent flow rates were also noted on the flow sheet. When
more than one feedstock was analyzed, stream compositions
were included for each feed.
In the process flow diagram, the general direction of flow
is uniformly from the feedstock and ingredient supplies on the
left to the product fuel on the right-hand side, with residuals
shown along the bottom of the diagram (represented by inverted
trapezoids). The flow path of the main process is always in-
dicated by a bold line, with thin lines indicating secondary pro-
cesses. A dashed line separates that portion of the process
specified by the process developer from the portion suggested
by the research teams for these analyses. In .several cases,
significant portions of the flow sheet were additions to the ele-
ments available in the literature or directly from the developer.
Consequently these flow sheets must be viewed as schematic
and not as representations of engineered systems. Sloping
11-20
-------
rectangles denote pollutant cleanup units external to the main
process, and rhombic-shaped units represent intermediate
products and uses for which the source and distribution are not
shown on the diagram.
Temperatures, pressures, .flow rates, and sulfur contents
are noted on each flow sheet where this information was available.
More complete stream compositions and sulfur: and energy bal-
ances for the overall process are reported in tables accompanying
each flow diagram. All physical quantities indicated directly on
the flow sheet or in the accompanying tables are reported in both
metric and English units, shown in Table 11-11.
Table 11-11
Units Used in Process Display
Quantity
Mass
Volume (liquid)
Volume (gas)
Length
Pressure
Temperature
Energy
Metric Unit
Ton (1000 kg mass)
Liter
o
Cubic meter, m°
Meter, m
Kilogram (force)/square
centimeter, kg/cm2
Degrees Celsius, °C
Kilocalorie, kcal
English Unit
Ton (2000 Ib mass)
Gallon, barrel
Cubic foot, ft3
Foot, ft
Pound (force)/square
inch (gauge and absolute),
psig and psia
Degree Fahrenheit, °F
British thermal unit, Btu
The assumed production rate, feedstock composition, pro-
cess variations, and any other specific characteristics are indi-
cated in explanatory notes accompanying each process display.
The approach used in obtaining the information presented
in the flow sheets and balance tables given in this section of each
synopsis and in the following sections on pollution control in-
volved:
11-21
-------
Literature review
Discussions with process developers
Engineering judgments and estimations.
It was found repeatedly that published information on the processes
studied was either incomplete or inconsistent. In these cases,
process developers or licensors were contacted directly to ob-
tain additional information. Although those contacted were
typically very cooperative, it was not always possible to obtain
the required data. In these cases, the research teams were
obliged to include cleanup processes and/or to develop estimates
of physical quantities based on experience with similar systems
and applying informed engineering judgment and estimating tech-
niques. The data presented in this report, though not always in
agreement with that published elsewhere, represents the most
up-to-date, accurate information available at this time (1973).
All assumptions which have been made in developing a complete
description of the pollution control system for each process
have been stated explicitly and the portion of the process speci-
fied by the developer is always clearly indicated.
(3) Process Description
Brief descriptions of each major process operation are
included in each process summary to indicate explicitly how
the process feed is converted to a desulfurized fuel and how the
major pollutant streams are separated, treated, recovered and
disposed. The specific steps defined for each process are de-
scribed and are keyed to the flow sheet given earlier. Sources
of all pollutant streams are shown on the flow sheet and described
in this section. A more complete discussion of the sulfur bal-
ance and total thermal energy balance calculated earlier is also
included in each summary.
(4) Discussion of Pollution Control Processes
Each process summary contains a discussion of control
processes or proposed systems to handle effluent streams gen-
erated or anticipated. Possible alternatives to these control
systems are also presented. Procedures for disposal and/or
utilization of process by-products in environmentally acceptable
11-22
-------
ways are also described in this section of the synopsis. The
generic pollution control unit processes selected are not dis-
cussed in detail in each clean-fuel process but instead are pre-
sented separately in Chapter III of this report. The feasibility
of using alternate processes to minimize emissions is discussed
in detail in this comprehensive review which is thus cited in
each process summary. The approach used to obtain information
and specific data on pollution control processes was similar to
that used in analyzing the clean-fuel processes themselves. Data
not available in the literature was obtained directly from process
developers or was developed during the course of this study by the
research teams assigned to each process. Each of the major
waste streams previously identified in the Process Description
section is discussed, and the control schemes applied to treat
each stream are evaluated. Fuel waste streams, solid residues,
and other pollutants escaping treatment are described, quantified
and possible disposition suggested. The intent of this section of
the process synopsis is to present sufficient detail to permit an
assessment to be made of the potential pollution control approach
which can be used and the problems which must be solved; how-
ever, as noted earlier, a complete optimization of the pollution
control system for each process and each feedstock was well
beyond the scope or intent of this study.
(5) Costs of Pollution Controls
In the concluding section of each process analysis, the
economic costs associated with the cleanup unit processes se-
lected for each scheme studied are calculated and presented.
Two standard accounting procedures (Discounted Cash Flow and
Utility Financing) were used in each case. The methodology
for performing the cost analysis was developed at the beginning
of the study and then applied uniformly to each process analyzed.
The procedures for determining the costs of pollution control are
described in detail in Chapter III of the report and are not re-
peated in each process synopsis.
II-23
-------
The principal elements of the approach used in meeting the ob-
jectives of the research program have been described in this chapter.
A case study approach was adopted in order to permit a reasonably
detailed technical analysis of the pollution control problems associated
with representative clean-fuels processes to be developed. The
synopses which were prepared for each process selected are presented
in Section 2 of this report. The following chapter includes discussions
of several technical topics which are common to all the processes con-
sidered and serves as a reference for each synopsis.
11-24
-------
III. DISCUSSION OF SELECTED TOPICS COMMON
TO ALL PROCESS DESCRIPTIONS
There are several topics which are pertinent to all of the clean
fuel process descriptions presented in Section 2 of this report. They
include:
Characterization of the pollutant streams emitted by the
processes
Description of the alternative processes available for
pollution control
Analysis of methods for integrating cleanup processes
into the primary flow
Development of the cost analysis methods used to compute
the cost of pollution control for each process.
Rather than repeat essentially the same information for each of the
processes considered, these topics are discussed in this chapter and
reference is made to specific processes where appropriate. The
process descriptions contain implicit reference to, and draw exten-
sively from, the material presented here.
1. CHARACTERIZATION OF DISCHARGES
The development of emission standards for clean-fuel processes
requires a detailed characterization of discharges from these pro-
cesses, as well as a discussion of the effectiveness and costs of
applicable control processes. The first of these topics is discussed
in this section under the following headings:
Sources
Waste streams
Sulfur and nitrogen emissions
Trace elements
JJJI-I
-------
(1) Sources
The sources of significant waste streams from clean fuel
processes are presented in Table III-l. For each source, the
table gives the associated waste streams, their principal com-
ponents, and manner of their disposition.
The sources of waste streams are many and varied. In
fact, there is hardly a step in any of the clean-fuel processes
that does not represent a potential source of environmental
pollution. In designing the process, it is essential to recognize
and characterize these sources, and then, to channel their waste
streams to recover valuable materials and to minimize environ-
mental impact.
The sources of waste streams can be grouped in four
categories:
Storage and preparation of feedstocks and other
supplies
Clean-fuel process operations
• Waste stream control and treatment processes
Ancillary activities.
The first category comprises storage of coal and shale
stockpiles, coal and shale cleaning and crushing operations,
treatment of water supply to remove dissolved salts that would
deposit in boilers or affect the process adversely, and genera-
tion of additional process supplies such as steam, oxygen, or
hydrogen. These sources and their discharges are rather easy
to characterize because the associated operations typically are
well-known, and the pollution problems are relatively easy to
correct.
On the other hand, the sources of discharges from clean-
fuel process operations are very difficult to characterize because
the commercial, "real world" versions of most of these processes
exist only in reports of engineering feasibility and design studies.
These reports, in turn, consist largely of extrapolations from
process development and pilot plant experience, where the main
III-2
-------
Table III-1
Nature and Sources of Major Waste Streams
Waste Stream
Coal and shale dust
Stack gas
Acid gas
Exhaust emissions
Steam
Runoff/ teachings
Wastewater
Slowdown
Cooling waler
Sanitary sewage
Rocks and debris
Tar
Char
Ash
Spent shale
Spent catalysts .
Spent purifying media
Sludge
Refuse
Principal Components
Carbon particles
See Table UI-2
See Table 1IIO
Nitrogen
See Table 111-2
See Table 111-2
CO, HC, NOX, particulates
Pyrites, sediment, oil.
organic matter
Dissolved salts
See Table III-3
See Table 11 1 -3
See Table HI -3
Dissolved salts
Heat
Organic/nitrogen compounds
See Table I1I-4
See Table III-4
See Table III -4
See Table 1114
Cobalt, molybdenum, iron
Bauxite
Nickel
Iron or zinc oxide
Activated carbon
Sulfites
Sources
Feedstock crushings & grinding
Steam generation
Stack gas cleanup
Oxygen generation
Product of shift conversion
Wastewater treatment
Automobile traffic
Cooling towers
Feedstock storage
Ash ponds
Land surface
Water, treatment
Gasification
Product of shift conversion
Methanation
Boilers/cooling towers
Heat exchangers
Washrooms
Feedstock cleaning
Gasification
Gasification
Gasification
Steam generation
Gasification
Shift conversion
Sulfur recovery
Methanation
Final purification
Final purification
Stack gas cleanup
Wastewater treatment
Work areas
Disposition
Containment and recovery
Tre atmen t/discharge
Discharge
Discharge
Sulfur recovery
Sulfur recovery
Discharge
Discharge
Treatment/utilization
Tre atmen t/utilization
Discharge
Treatment/utilization
Treatment and reuse
Treatment and reuse
Treatment and reuse
Treatment and reuse
Utilization/discharge
Treatment and disposal
Disposal
Utilization
Utilization
Disposal/ utilization
Disposal/utiliza tion
Disposal
Treatment and recovery-
Treatment and disposal
Treatment and disposal
Treatment and disposal
Disposal
Treatment and disposal
Treatment and disposal
Treatment and disposal
Process/Unit
Cyclone separators, bag filters, enclosure
^ Wellman-Lord, limestone injection, lime scrubbing, catalytic
/ oxidation, double alkali, Citrate
H2S can be selectively removed from a gas stream by a rectisol.
hot potassium carbonate, Sulfinol, MEA, DIP A or similar process
) Ctaus and Stretford processes recover elemental sulfur from H^S-rich streams
* Tail gases from these units can be treated by incineration, Beavon, Wellman-
Lord, or SCOT process
Neutralization
Chevron, Phenosolvan, Phosam, oxidation, biologict t
Waste heat recovery, cooling towers
Biological
Conversion to oij, direct gasification, direct combustion, indirect combustion
Gasification, desulfurization, direct combustion
Landfilling/incorporation in construction materials
Minefilling
Oxidation
Oxidation
Oxidation
Oxidation
Dewatering, oxidation, land filling
Incineration, landfilling
Ref.
-------
emphasis was on "making the process work"; comparatively little
attention was paid to potential sources of environmental pollution
due to the limited resources made available for these studies.
Pollution control and treatment processes comprise
principally those applicable to coal dust, stack gas, acid-gas,
wastewater, char and tar, and heat. Inasmuch as these proc-
esses are bound by laws of conservation of mass and energy,
their task is not to eliminate polluting substances but merely
to convert them into a less offensive form, chemical species,
or state of aggregation. These processes, however, are still
capable of generating undesirable discharges that may be treated
further or monitored to ensure compliance with applicable
standards.
Finally, sources associated with ancillary activities
catering to human sanitation and transportation needs are
probably easiest to deal with because they are so common, and
the solutions are so readily available. It is important, however,
they not be left unattended in the effort to meet the more pressing
requirement to control discharges from the clean-fuel process
operations.
(2) Waste Streams
The major gaseous, liquid and solid waste streams asso-
ciated with clean-fuel processes are listed in Tables III-2, 3
and 4. For each waste stream, these tables provide the principal
components, typical sources, manner of disposition, control or
treatment process. The particular processes are discussed in
considerable detail in the individual chapters of Section 2.
The principal airborne waste streams comprise dust from
handling and grinding of coal and shale, stack gas from the
steam plant, and acid-gas from the primary process stream
following the shift conversion operation. Coal dust can be con-
trolled rather effectively by cyclone separators and filter bags
and by enclosure of the feedstock preparation unit. Stack gas
from steam plants fired by fuel gas or low-sulfur fuel oil can
often be vented directly to the atmosphere without violating the
current power plant emission standards. Control of acid-gas
removal stages following CO shift conversion takes place in the
III-4
-------
much more complex and costly sulfur recovery stage. Eventually
the hydrogen sulfide is oxidized to free sulfur, which can be sold,
while the carbon dioxide is vented to the atmosphere.
The liquid waste streams consist chiefly of runoff and
leachings from coal and shale stockpiles, hot cooling water from
heat exchangers, and wastewater from unit processes. The
latter includes wastewater from gas scrubbers following the
gasification stage and condensed moisture of the gas streams
from the CO-shift conversion and methanation stages. Runoff
and leachings would receive only minimal treatment, primarily
to neutralize the acidic pyrites, before discharge; hot cooling
water would be used to supply heat to other process units.
However, proposed process wastewater treatment is again com-
plex and costly and leads to recovery of ammonia and phenols,
conversion of hydrogen sulfide to sulfur in the sulfur recovery
stage, and recycling of the treated water to the scrubbers or
cooling towers. A solid, inert sludge remains to be treated
and disposed.
The solid waste streams consist of mineral debris, spent
shale, ash, and sludge, as well as tars and chars. The former
would be disposed in a landfill or an abandoned mine, while tars
and chars can be used as fuel following treatment to reduce their
high-sulfur content.
The potential amounts of major discharges from a plant
producing 7. 1 x 106 m3/day (250 x 106 ft3/day) of pipeline gas
from Illinois No. 6 coal containing 3. 7% sulfur are listed below.
III-5
-------
Table III-2
Composition of Selected Gaseous Waste Streams
(Vol. Percent)
Source
Disposition
Process
Feedstock
Product
H,
CO
C02
CH4
C,+
COS
so2
H,S
N2
°2
H2O
Feedstock Preparation
Discharge
COED
Utah
Coal
Syn-
trade
13.98
70.52
15.50
CO2 Accep.
N. Dakota
Coal
SNG
0.17
22.06
0.06
0.02
0.02
0.01
54.44
0.19
23.03
Acid Cases
Sulfur Recovery
Synthane
Pittsburgh
Coal
SNC
0.1
93.6
1.5
4.8
Hygas
Montana
Coal
SNC
70.00
0.14
29.86
Hygas
Illinois
Coal
SNC
70.0
0.17
29.83
Lurgi
Navajo
Coal
SNC
0.75
0.41
95.45
0.58
0.52
1.05
0.54
0.70
COED
Utah
Coal
Syn -crude
1.30
0.77
89.32
0.34
0.39
1.88
6.00
COED
Illinois
Coal
Syn-crude
1.3
0.7
75.3
0.4
0.3
22.0
SRC
Kentucky
Coal
Coal
27.55
72.45
Texaco
Residual
Oil'
Synthesis gas
69.70
21.84
Texaco
Residual
Oil
Hydrogen
54.42
2.00
42.05
1.53
Discharge
Hygas
Montana
Coal
SNC
0.45
0.03
97.90
0.16
1.46
12 ppm
5 ppm
Texaco
Residual
Oil
Hydrogen
0.20
91.58
0.10
o.o:
5 ppm
8.10
Sulfur Recovery
Discharge
Lurgi
Navajo
Coal
SNC
0.3S
0.15
83.53
0.46
0.46
10.58
2.25
1.89
Steam Plant
Qeanun
Synthane
Pittsburgh
Coal"
SNC
17.3
. 0.13
78.4
2.9
1.3
S Recovery
SRC
Kentucky
Coal
Coal
15.99
2.27
79.00
2.74
Synthane
Pittsburgh
Coal
SNC
17.3
0.06
78.4
2.9
1.3
Texaco
Residual
Oil
Synthesis gas
11.72
0.02
74.65
2.59
11.02
Texaco
Residual
Oil
Hydrogen
12. 36
0.04
75.10
2.60
9.90
-------
Table III-3
Composition of Selected Liquid Waste Streams (mg/1)
Source
Disposition
Process
Feedstock
Product
COD
DS
SS
CO2
CH4
Phenol
NH3
Cyanide
Thiocyanate
H2S
Sulfur
Gas Scrubber
Treatment
Syn thane
Pittsburgh
coal
SNG S
19,000
23
1,700
11,000
0.6
188
Gas Scrubber
Treatment
Texaco
Residual
Oil
Acid-Gas Removal
Treatment
Synthane
Pittsburgh
Coal
'nthesis Gas or H2 SNG
5
5,000
10
300
25
13,000
300
Methanation
Treatment
Lurgi
Navajo
Coal
SNG
50
170
Gas
Scrubber
Treatment
Synthane
Illinois
Coal
SNG
15,000
18,900
600
2,600
8,100
0.6
152
Gas
Scrubber
Treatment
Synthane
Kentucky
Coal
SNG
19,000
55
3,700
10,000
0.5
200
-------
Table III-4
Composition of Selected Solid Waste Streams (wt. percent)
oo
Stream
Disposition
Process
Feedstock
Product
C
H
0
N
S
Ash
Minerals
Yield (%)
Char
Syn thane
Pittsburgh
Coal
SNG
71.4
0.9
1.8
0.5
1.5
23.9
30.5
Utilization
COED
Illinois
Coal
Syn-crude
73.4
1.3
0.3
1.2
3.1
20.7
52.9
COED
Utah
Coal
Syn-crude
85.2
1.1
1.0
1.4
0.7
10.6
46.5 .
Tar
Utilization
Synthane
Pittsburgh
Coal
SNG
77.4
7.0
7.3
—
5.5
2.8
3.5
Ash
Disposal/Utilization
Hygas
Montana
Coal
SNG
10.2
0.1
89.7
8.1
CO2 Accept.
No. Dakota
Coal
SNG
5.2
94.8
0.7
SRC
Kentucky
Coal
Coal
3.95
0.54
95.51
6.7
-------
The ranges in values are indicated by variations among dif-
ferent gasification processes and yield uncertainties. *
Sulfur
Oils and tars
Benzene
Ammonia
Phenols
Hydrogen cyanide
Mercury
350-500 tons/day
0-400 tons/day
50-300 tons/day
100-150 tons/day
10-70 tons/day
0-2 tons/day
5-10 Ibs/day
318-454 m tons/day
363 m tons/day
45-272 m tons/day
91-136 m tons/day
9-64 m tons/day
0-2 m tons/day
2-5 kg/day
(3) Sulfur and Nitrogen Emissions
The principal contaminant of fossil fuels and the chief
justification for clean-fuel processes are the compounds of
sulfur which are converted to sulfur dioxide (SO2) upon com-
bustion. A portion of sulfur dioxide in the atmosphere is
further oxidized to sulfur trioxide (803) which then combines
with moisture to form a sulfuric acid (H^SO^.) mist. This mist
corrodes exposed surfaces of buildings and other outdoor pro-
perty, destroys the vital leaf surfaces of plants, and is thought
to contribute considerably to the incidence of respiratory afflic-
tions in humans and animals.
The sulfur in coal may be present in either organic or
inorganic forms. The organic compounds consist primarily of
sulfides, thiophene, and benzothiophene derivatives and con-
stitute the principal sulfur fraction of low-sulfur coals. The
inorganic sulfur is present principally as iron disulfide (FeS2)
in its mineral configurations of pyrite of marcasite, though
lesser amounts appear as sulfates (CaSO4 or MgSO4). Pyritic
sulfur is generally easier to remove than organic sulfur and is
the major sulfur constituent of high sulfur coals.
"Final Report, The Supply Technical Advisory Task Force -
Synthetic Gas Coal," Federal Power Commission, April 1973.
III-9
-------
The distribution of sulfur among the outputs of selected
clean-fuel processes is reported in Table III-5 for various feed-
stocks, in terras of weight percent of each feedstock. It will
be noted that in the case of coal gasification processes, over
75 percent of the original sulfur is recovered in elemental form.
In general, the fraction recovered increases with the original
sulfur content of the feedstock.
The next major contaminant of fossil fuels, after sulfur,
is nitrogen which forms nitrogen oxides (NOX) upon combustion,
and ammonia (NHg) upon hydrogenation. Ammonia can be further
converted to toxic hydrogen cyanide (HCN) by reaction with
hot coal. Nitrogen occurs primarily in the form of organic com-
pounds such as pyrroles and pyridines. Its content in coals varies
between one and two weight percent.
(4) Trace Elements
Trace elements are usually defined as those present in the
earth's crust in amounts of less than 1000 ppm. Nearly all trace
elements exhibit an enrichment in fossil fuels relative to their
respective average crustal abundance, probably because they were
accumulated by the vegetation from which the fuels were formed.
Of the approximately 40 trace elements found in coal and oil,
seven have been labeled hazardous because of their toxicity and
volatility:
Arsenic • Lead
Beryllium • Mercury
• • Cadmium • Selenium
Fluorine
Little'is known at this time about the fate and distribution
of these elements in clean-fuel processes. The EPA is currently
evaluating this problem in a separate experimental program. The
final results have not yet been published but the initial data is
presented in document EPA-650/2-73-004, August 1973.
(5) Heat
The conversion of feedstocks to clean fuels and other useful
products invariably involves some loss of energy. This energy
111-10
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Table III-5*
Sulfur Distribution Among Outputs of Clean-Fuel Processes
(Weight Percent of Feedstock)
--~^^^ Process
-~-^Feedstock
Carrier "^ — ^^^
Coal
Oil
Tar
Char
Ash
Spent Mineral
Sulfur
Sulfur Dioxide
Sul fides
Sulfates
Total
Synthane
Pittsburgh Illinois
' Coal Coal
0.19 0.04 '
1.19 2.65
0.20 0.18
0.02 0.03
1.60 2.90
Hygas
Montana Illinois
Coal Coal
0.007 0.005
0.005 0.03
0.48 3.86
( 0.01 J0.03
0.51 3.93
CO-) Acceptor
North Dakota
Coal
0.45
0.09
0.006
0.04
0.59
Lurgi
Navajo
Coal
0.005
0.01
0.66
0.01
0.005
0.69.
COED
Utah Illinois
Coal Coal
0.005 0.03
0.33 0.47
0.14 2.80
0.03
0.47 3.33
SRC
Kentucky
Coal
0.64
0.01
0.04
2.43 ,
0.02
3.14
Gas Combustion Retort
Colorado
Shale
0.48
0.10
0.02
0.60
Texaco Partial Oxidation
(Synthesis) (Hydrogen)
Residual Fuel Oil
6.1C 6.07
0.10 0.13
6.20 6.20
•Ultimate disposition of the feedstock sulfur is presented in Table 1-3. It is presented there as a
percentage of the input sulfur.
-------
loss can be ascribed to the combined values of heats of combus-
tion of waste streams, sensible heatsv of cooling water and product
and waste streams, and heat lost to the atmosphere, either through
direct radiation or through discharge of heated air or evaporation
of cooling water.
The distribution of the heat of combustion of the feedstocks
among the product and waste streams of the various processes
is presented in Table III-6. For each process, the table lists
the feedstocks, the product, and the percentage of feedstock
energy retained by the various product and waste streams. . The
last row reports the combined heat losses through airborne and
waterborne discharges. The specific calculations are shown
under each process synopsis in Section 2 of this report.
It will be noted that the heat losses vary between 16 and
34 percent, while the fraction of heat of combustion retained in
the clean fuels, also referred to as the energy conversion effi-
ciency, fluctuates rather widely, depending on which products
are regarded as "clean fuels. "
In concluding this section on pollutant discharges, it is appro-
priate to compare the discharge from a typical coal gasification pro-
cess with those from a coal-fired electric power plant yielding the .
same quantity of product energy. Table III-7 makes this comparison.
The sulfur emissions from a power plant meeting EPA new source
performance standards for SC>2 are 30 times the total sulfur emissions
from the coal gasification facility. Additionally, the power plant re-
quires more coal, generates more solid waste and has higher thermal
emissions.
2. ALTERNATIVE PROCESSES FOR POLLUTION CONTROL
The potential emissions generated in processes for clean-fuel
production may be classified according to type:
Noise . Liquids
Thermal . Gases
Solids
Sensible heat is defined here as the heat given up by a substance
in reducing its temperature to 16 degrees C (60 degrees F).
111-12
-------
Table III-6
Distribution of Energy in
Clean-Fuel Products
(Percent of Feedstock Energy)
Process
Feedstock
Product
Gas
Oil
Coal/coke
Char
Tar
Ash
Sulfur
Ammonia
Phenol
Other products
Heat recoverec
Cooling water
Other heat*
Heat losst
Syn thane
Pittsburgh
Coal
SNG
64.0
4.:
0.3
0.5
64.0
25.8
5.2
31.0
Hygas
Montana Illinois
Coal Coal
SNG SNG
60.1 60.4
6.5 3.7
0.2 1.3
0.2 0.6
0.1 0.1
17.1 66.1
13.9 13.9
19.0 20.0
32.9 33.9
C02
Acceptor
North
Dakota
Coal
SNG
59.2
0.3
0.6
10.4
70.5
9.5
20.0
29.5
High-Btu
Lurgi
Navajo
Coal
SNG
56.2
5.1
6.6
1.6
0.3
5.1
0.8
75.7
5.9
18.4
24.3
COED (gas reform.)
Utah Illinois
Coal Coal
Syn-crude Syn-crude
4.4 3.8
31.8 23.0
47.5 51.3
0.04 1 .0
0.16 0.2
83.9 79.3
16.1 20.7
16.1 20.7
SRC
Kentucky
Coal
Coal
7.4
67.2
0.8
0.4
1.0
76.8
21.6
1.6
23.2
Texaco
Partial Oxidation
Residual Oil
Synthesis
Gas Hydrogen
79.9 69.3
1.3 1.1
3.5
81.2 73.9
6.3 5.4
12.5 20.7
18.8 26.1
Gas Combustion
Retort
Mahogany
Oil Shale
Syn-crude
74.7
6.5
0.2
0.7
82.1
17.9
17.9
Low-Btu
Lurgi
Navajo
Coal
Utility
Gas
66.5
4.1
7.4
0.3
1.0
0.8
3.8
83.9
16.1
16.1
U-Gas
Pittsburgh Coal
Combined
Cycle Utility
Fuel Gas Gas
73.2 79.3
4.4 4.1
1.3 1.3
6.6 3.3
85.5 88.0
14.5 12.0
14.5 12.0
Koppers-Totzek
Dlinois Eastern Western
Coal Coal Coal
Utility Utility Utility
Gas Gas Gas
69.8 74.6 73.2
1.7 0.3 0.2
4.5 0.9 1.2
76.0 75.8 74.6
20.0 19.4 19.7
4.0 4.8 5.7
24.0 24.2 25.4
* Obtained by difference: includes sensible heat of product and waste streams, heat of combustion of waste, streams, and heat lost directly to the atmosphere.
t Defined as the percent of heat of combustion of the feedstocks that is lost through heat of combustion of waste streams, sensible heat of cooling water and
product and waste streams, and heat lost directly to the atmosphere.
-------
Table III-7
Comparison of 250 Billion Btu/day Gas from Coal
with 3000 MW Electricity from Coal
Delivered Energy
Gas
Electricity
Product, Btu/day (kcal/day)
By-Products, Btu/day (kcal/day)
Coal Consumption, short ton/day (m ton/day)
Thermal Emissions, Btu/day (kcal/day
Sulfur Emissions, short ton/day (m ton/day)
S02
H2S
COS
Total sulfur losses (as S)
Waste, short ton/day (m ton/day)
Ash (dry)
Gypsum (dry)
250 x 109(63x 109)
24 x 109(6x 109)
17,517(15,891)
140 x 109(35x 109)
5.6(5.1)
0.12(0.11)
5.55 (5.03)
6.05 (5.49)
2248 (2039) .
246 x 109(62x 109)
27,400 (24,857)
401 x 109(101 x 109)
388 (352)
194(176)
3520(3193)
5150(4672)
Basis: Illinois No. 6 coal, 4.2 percent sulfur, 11,800 Btu/lb (6556 kcal/kg), 6-1/2 percent moisture;
electricity generation at 38 percent efficiency, 1.2 Ib SO2/106 Btu, (2.2 kg SO2/106 kcal);
gas manufacture by the Hygas process.
-------
The alternative processes available for control of these emissions
are discussed in this section. Principal emphasis in the study was de-
voted to control of gaseous waste streams, particularly those containing
sulfur. Hence the cleanup processes pertinent to these streams receive
greatest attention.
(1) Noise
Noise pollution was not specifically studied in this pro-
gram. The noise generated in these processes will be similar
to that found in most chemical processes as generated by pumps,
compressors, pressure relief valves, and so forth. Standard
engineering practices can be used to control noise emissions.
(2) Thermal
The level of thermal pollution is inversely proportional
to the process efficiency attained. The heat content of the pro-
ducts of the process is always less than that of the feedstock
because energy was required for extraction and purification.
This energy difference appears as heat lost through thermal
discharges. In general, the efficiencies of these processes are
quite high as compared with other forms of energy transforma-
tion. Typically, the coal gasification processes have overall
efficiencies, including by-product credits, on the order of 55 to
70 percent, depending upon the process and the coal feedstock
used. Therefore, the typical thermal discharges are on the
order of 0.4 to 0. 8 units of energy for each unit of clean energy
produced. This may be compared with a modern nuclear power
plant which operates at an efficiency of 30 to 33 percent, dis-
charging two units of energy per unit of output. Fossil-fueled
power plants are somewhat more efficient than nuclear plants
(to 38 percent) but still produce a thermal discharge of 1.6 units
per unit of output, double the emissions from the worst gasifi-
cation facility.
The disposition of thermal discharge between air- and
evaporative-cooling is typically not specified by process
developers. Only the Lurgi gasification process, as designed
for the El Paso Natural Gas Company, presents sufficient
engineering data to evaluate evaporative losses. In this case,
extensive use is made of air coolers, and less than 50 percent
III-l5
-------
of the thermal discharge is by evaporative cooling. The data
presented for the Hygas process indicate a similarly low fraction
of evaporative cooling, but this process has not yet been as
thoroughly engineered as the El Paso project .
(3) Solids
The solid discharges from clean energy conversion proc-
esses may include the following:
Sulfur
Spent catalysts
Gypsum
Pyrites and dust
• Ash and rock . .
Char
Though not all of these discharges will be produced in every
clean-fuel process, each possible pollutant is discussed below.
1. Sulfur
A major function of most of the clean-fuel conversion
processes is the elimination of sulfur from the feedstock.
Usually, this will appear in the form of elemental sulfur as
a by-product from the process. This solid discharge is the
most environmentally attractive chemical form for the
discharge of sulfur. Although the literature contains
references to aerobic and anaerobic bacterial action on
elemental sulfur, it is still the preferred form of dis-
charge because it is relatively inert, occupies a relatively
small volume, and is potentially saleable. It is expected
that the market for elemental sulfur will continue to expand
for the production of fertilizers, and that this by-product
can be sold if competitively priced. The long-term storage
capabilities of block sulfur are amply proved by the Frasch
blocks in the southern United States. Alternatively, the
sulfur may be disposed of as clean landfill or returned to
the mine.
111-16
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Most of the solid-based processes for the manu-
facture of clean energy will require substantial storage
facilities for the feedstock. In general, a 30- to 60-day
supply will be maintained to ensure continuity of plant
operation. In general, this stockpile should not cause
problems except, perhaps, in the case of some high-
volatile subbituminous and lignite coals which are subject
to spontaneous combustion. This problem has also been
faced by the electric utility industry. The measures used
to ensure plant safety, such as stockpile thermal monitoring,
stockpile rotation, and water sprays, should also prove
effective for the control of any sulfur emissions resulting
from spontaneous combustion.
2. Spent Catalysts
A wide variety of catalysts may be used in the vari-
ous clean energy conversion processes. These catalysts,
whether used to promote a chemical reaction or used as
final purification sorbents, deactivate with time or during
plant upsets. Generally, catalysts must be replaced on
a schedule of one or two years, although sorbents may
require more frequent replacement. Some sorbents, such
as metal-activated carbon or iron-oxide sponge will even-
tually become contaminated with elemental sulfur and,
after possible pyrophoric tendencies have been eliminated
by deactivation, may be discarded with the ash from the
process. In most cases, however, the catalysts and sor-
bents will be returned to the supplier for reclamation.
This study did not consider the regeneration techniques
employed by the supplier for each of the various potential
catalysts which might be used in clean-fuel conversion
processes.
3. Gypsum
Most of the processes analyzed require steam and
shaft horsepower that can be produced by burning a por-
tion of the feedstock. In this case, the boiler stack gas
must be desulfurized. Many stack-gas cleanup processes
are currently under development; wet-lime scrubbing is
111-17
-------
now considered by the EPA to be a proved process for this
application. The primary by-product from this operation
is gypsum, an inert solid which is composed of hydrated
calcium sulfate. This gypsum will leave the process, to-
gether with unreacted limestone, as a thixotropic sludge
(a gel becoming a liquid when agitated) which will pose
handling problems. It will probably be disposed of with
the ash from the system to either landfill, sludge drying
beds, or minefill. The ash may stabilize the thixotropic
character of the sludge.
4. Pyrites (Sulfides, e.g., FeS2)
Many of the processes that operate on coal may
function more efficiently if the coal is prewashed to re-
duce its sulfur content through pyrite separation. This
operation will increase the quantity of product which can
be made from the coal* but will reduce the amount of
elemental sulfur by-product. The pyrites, which are
eliminated from the coal in the washing operation, pose
special disposal problems. When exposed to long-term
action of air and surface water, the pyrites will decompose,
forming ferrous sulfide and acidic liquor as characterized
by acid-mine water. A satisfactory means for disposal
of these pyrites is blending them with excess quantities of
limestone which would neutralize any acid formed by
weathering.
5. Dust
Any of the processes which use a solid feedstock
will generate dust during crushing and grinding opera-
tions. The techniques for controlling this dust (such as
negative pressure grinding, cyclones, baghouses, and
spray rinses) are well known in the electric utility indus-
•try where grinding for cyclone furnaces has been practiced
for many years.
Hydrogen needed for clean-fuel production is not lost to H S
formation.
111-18
-------
•S
6. Ash and Rock
The solid residues which were present in the feed-
stock for these processes are generally considered to be
ash if they have very little heating value, as in the case
of the Hygas, Lurgi, and Texaco Partial Oxidation pro-
cesses. When large quantities of inert substances are
present, as in the Gas Combustion Retort process, these
residues are considered rock. On the other hand, when
the carbon removal is not as complete, as exemplified by
the Synthane or COED processes, the by-product is called
char.
The ash and rock are relatively inert. Any sulfur
remaining in the ash will be fixed into the carbon lattice
as a stable material. The ash may have by-product value;
for example, as concrete aggregate, paving material, or
as an ingredient in cement block manufacture. More often
the ash will be disposed of as landfill or perhaps by re-
turning to the mine.
Rock from the Gas Combustion Retort process for
producing shale oil poses special problems. Large ton-
nages of oil-bearing shale must be processed to extract
reasonable quantities of product, and the kerogen-depleted
rock will not compact into the same volume as the original
shale. In this case, the waste-product rock will more
than fill the mine from which it was taken. It is assumed
that, in the mountainous area where shale is found, remote
canyons can be found for rock disposal. Aesthetics can
be maintained by subsequent grading and landscaping
efforts.
7. Char
The char from some of the clean-fuel manufacturing
processes may contain significant carbon due to process
operating conditions. The carbon should be recovered in
the form of energy, but the residual sulfur must still be
removed for environmental reasons. Recovery of the
energy in char may also be necessary for the overall
economy of the process. Substantial energy may be
III-19
-------
required to make process steam or to operate an air
separation plant for oxygen production. The overall
efficiency of most processes is improved if the waste
char is used to provide that energy.
Sulfur can be removed from the char either before,
during, or after combustion. In this respect, the com-
bustion of high-sulfur char is similar to the combustion
of high-sulfur coal. Therefore, desulfurization efforts
toward the clean combustion of coal are applicable, with
modifications, to the environmentally acceptable utilization
of char; however, the sulfur in the chars from these proc-
esses has not been characterized as to quantity or chemical
composition since these are functions of the process oper-
ating conditions and the characteristics of the initial feed-
stock. Different sulfur characteristics in the char may
eliminate some of the coal pretreatment processes cur-
rently under consideration; for example, the TRW Leaching
process and the IGT Flash Desulfurization process. How-
ever, various chars may be amenable to new types of pre-
treatment because of the changed characteristics of the
sulfur. Many of the process developers have been studying
the problem of sulfur removal from char, as evidenced by
the char desulfurization concept described in the COED
process synopsis (Chapter VIII). The following discussion,
however, is limited to sulfur removal processes that have
been proved or are currently under advanced development.
The processes for sulfur removal during and after
combustion of char would be quite similar to those under
development for coal; for example, the sulfur may be re-
moved by combustion in a bed of lime. Many processes
are being developed to desulfurize stack gas from the com-
bustion of coal; limestone injection, lime scrubbing,
catalytic oxidation, Wellman-Lord, double-alkalai, and
citrate processes, among others, are promising. Accord-
ing to the EPA, flue gas desulfurization is now a proven
process for stack-gas desulfurization from the combustion
of high-sulfur coal; it is assumed to also be satisfactory
for char combustion.
The subject of char precombustion treatment also
deserves attention. As mentioned above, many coal-
treating processes may not be applicable to char; however,
111-20
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chars can be gasified completely, making a fuel gas which
can then be desulfurized and utilized. This approach was
used by the El Paso Natural Gas Company in its applica-
tion of the Lurgi process. In this case, fuel gas desul-
furization used the same general acid-gas cleanup tech-
nique, including the same regenerator, that was employed
for the primary gas-treating system. A Lurgi gasifier,
using air as the oxidant, could, therefore, be used on most
chars to produce a low-Btu fuel gas (approximately
200 Btu/ft3 (1780 kcal/m^), and a relatively clean ash
which could easily be disposed.
The Hygas flow sheet depicts a low-Btu gas system
(U-Gas) for steam production. Either the steam-oxygen
gasifier of the Hygas system or the steam-air gasifier
of the U-Gas system would be suitable for converting
chars into consumable fuel gas and disposal ash.
A Koppers-Totzek gasifier would also be applicable for
this purpose. It operates at high temperatures so that
all of the carbon and sulfur would be gasified, leaving a
sulfur-free ash for disposal. Another alternative gas-
ification technique would be the use -of the bottom stage of
the Bi-Gas reactor for the production of a low-Btu gas
from char. Here again, the high-temperature slagging
Bi-Gas reactor would consume all of the carbon and sulfur
in the char, producing a clean ash. This unit of the Bi-
Gas reactor could also operate on air rather then oxygen
to produce a nitrogen-diluted fuel gas with low Btu con-
tent.
The COED process poses a special problem on han-
dling chars. The quantity of char produced from the COED
process contains much more fuel value than the energy re-
quirements of that process. The process developers, the
FMC Corporation, propose a lime-treatment system which
will result in a low-sulfur solid fuel product; however, the
volatile content of this char is low, so that it cannot be
•readily combusted. The COED system could be used in
combination with a variety of other processes which could
utilize the char. For example, the system could be operated
in conjunction with an electric utility, the COED char could
be gasified, and the resulting low-Btu fuel could be used
either under boilers or as feed for gas turbines. Using the
COGAS process developed by FMC, the char can be
111-21
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gasified to produce methane. The COED char could also
be steam-oxygen gasified and purified to make a synthesis
gas, feeding ammonia manufacture, methanol fuel produc-
tion, or fuel oil production by the Fischer-Tropsch reaction.
Of course, all of these options were available for the initial
coal feed, but the COED process extracts a liquid hydro-
carbon product initially, an option that may be attractive
for some process developers.
(4) Liquids
The liquid waste streams discussed will be categorized
into aqueous or organic streams.
1. Liquid Waste Disposal—Aqueous Streams
Future clean-fuel production plants will operate
much like modern oil refineries, and the wastewater dis-
posal problems will also be similar. The primary waste -
water streams for most of these systems will also resemble
the quench water from coke ovens.
The term "wastewater," as applied to these facilities,
is a misnomer. Although aqueous streams are generated
in the processes, these streams can be treated and re-
turned to the cooling water circuits, according to most
process developers. For example, the cooling water re-
quirements for a 250 million ft^/day (7.1 x 10°m3/day)
synthetic gas plant are approximately the same as those
for a 600 megawatt (MW) power station. As in a modern
power plant, the water is recirculated through cooling towers
to near extinction. Perhaps a small sidestream would be
withdrawn to evaporation ponds to eliminate a buildup of
soluble salts, but there will probably not be an aqueous
waste effluent to the local watershed from a modern, clean-
'fuel processing plant.
Contaminated aqueous streams are generated in the
clean-fuel processes because product or effluent gases
are either water-washed to remove soluble contaminants
or are quenched for temperature control. A third poten-
tial source of an aqueous stream is from simple conden-
sation of excess steam which might have been used in a
HI-22
-------
gasification process. The type and quantity of contaminants
found in the washwater are function's of the chemical
characteristics of the initial feedstock to the process, the
type and severity of treatment in the primary clean-fuel
unit process, and the type of subsidiary treatment upstream
of the water-wash, quench, or cooling-condensation steps.
In general, contaminants of the following types
might be found in an aqueous stream produced in a clean-
fuel process:
Particulate Matter. Fly ash and/or coal dust
may be found in the aqueous stream, partic-
ularly if a hot gas is quenched directly after
gasification. A prime example of this effect
is found with the Koppers-Totzek gasifier
which treats pulverized coal in a dilute-phase
reaction. Solids might also be found in the
overhead streams from other dilute-phase
gasifiers or fluid-bed operations using fine
coal feed. Similarly, partial-oxidation of
residua can produce a fly ash or soot.
Emulsified Tars and Oils. Many of the gas -
ification processes produce tars and oils in
varying amounts. They might emulsify into
the aqueous stream if the gasifier effluent is
directly quenched.
• Sulfurous Compounds. Most.of the gas streams
which are washed will contain sulfurous com-
pounds (primarily hydrogen sulfide and car-
bonyl sulfide) which are water soluble. A
small fraction of the sulfur in the gas will,
therefore, be dissolved into the wash water.
If ammonia is also present in the gas, this
sulfur solubility increases.
• Nitrogeneous Compounds. If the initial feed-
stock to the system contains chemically com-
bined nitrogen, as in the case of most coals and
oil shale, the nitrogen will partially hydro-
genate to ammonia in most clean-fuel processes
unless the primary reaction occurred at very
111-23
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high temperatures. The ammonia will be
removed from the gas stream in the water
wash.
Ammonia can react with carbon at high tem-
peratures to form hydrogen cyanide accord-
ing to the H-C-N thermodynamic equilibrium,
and in fact, cyanide is reported in many gas-
ifier effluents. This cyanide will dissolve in
the aqueous stream if the gasifier effluent is
quenched directly. It may also catalytically
hydrogenate to ammonia over a shift catalyst
if condensation and wash occur later in the
processing scheme.
Some process developers report nearly quan-
titative conversion of hydrogen cyanide with
hydrogen sulfide to form thiocyanates. Ther-
modynamics do not favor this reaction in a
reducing atmosphere, but the conversion is
rapid when the foul water is exposed to air.
• Phenolic Compounds. Aromatic-based feed-
stocks, such as coal, may produce small
quantities of phenols and cresols during the
primary treatment step if the processing con-
ditions are not severe. These hydroxylated
compounds are water soluble and will appear
in the aqueous stream.
These aqueous streams are purified in the clean
energy processes. The first processing step for foul
water streams in most cases is pressure reduction by
flashing. Much of the dissolved sulfur and carbon dioxide,
together with desirable gas constituents, are released
from solution in this step. The scavenged gas is usually
repressurized and placed back into the primary process
flow stream.
The next processing step involves solids removal.
Conventional solids-liquid separation equipment, such as
thickeners, clarifiers, and filters, can be applied. The
API separator, a device that will remove floating oil
111-24
-------
and insoluble solids simultaneously from the liquid, will
probably find wide application on streams that contain
insoluble matter. The solids from this system can be
blended with the ash from the system for. disposal, and
any floating oil marketed with that by-product (as discussed
later).
The remaining constituents in the foul water, with
perhaps the exception of thiocyanates and ammonia, are
subject to biological oxidation; however, processes are
available for their recovery. The economics of the
individual operation, with specific emphasis on the partic-
ular feedstock and primary process operation, will
determine whether the by-products will be recovered or
destroyed.
Phenols and cresols can be recovered by liquid-
liquid extraction techniques, as is indicated in the Lurgi
and Hygas process descriptions. The Lurgi process
makes large quantities of phenols which might pay for the
recovery process. The Hygas processing conditions, on
the other hand, do not favor phenol production; phenol and
oil recovery are employed simply to minimize the bio-
logical treatment of the final wastewater. The phenols,
together with the extractable oils, have a high by-product
value. For the process synopses presented in Section 2,
a. value of 4-1/4^ /lb was used for by-product phenol credit.
This price is approximately one-half of the present market
price for phenol because the by-product will require further
upgrading for oil and sulfur removal by a cooperating
refinery or specialty chemical manufacturer.
Ammonia is removed from the system by steam
stripping. Processes, such as the United States Steel
Engineers' Phosam process, are available to remove
ammonia selectively from the stripped gas; the other
constituents in the gas can be returned to the process.
'Alternatively, the ammonia can be recovered as a solu-
tion by selective stripping. Cyanide contamination should not
be expected if it were quantitatively converted to thiocyanate
by upstream oxidation in the presence of hydrogen sulfide.
The product ammonia (either anhydrous from the Phosam
process or in solution from selective stripping) should be
a readily marketable by-product.
111-25
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Wastewater after ammonia stripping will contain
small quantities (in the order of several ppm) of sulfur,
carbon dioxide, and cyanide. If phenol recovery were
practiced, the wastewater would contain very little of
these materials but will, however, contain most of the
thiocyanates which might have been formed during the
treatment processes. In general, this''stream will be
sufficiently pure to be a satisfactory input to the cooling-
water circuit of the facility.
To further treat the wastewater stream and the
effluents from boiler blowdowns, cooling tower blowdowns,
and plant sanitary facilities, it is expected that many clean -
fuel processes will employ biological oxidation.
2. Liquid Waste Disposal—Organic Streams
Many of the clean-fuels processes produce tars,
oils, and BTX (benezene-toluene-xylene) by-product
streams. The tars and oils have product value as fuel,
but the BTX stream has increased value as a component
of motor fuel due to its high octane rating.
A BTX stream is shown for both the Lurgi and Hygas
processes. It is removed from the product gas in the first
stage of treatment in a solvent-based acid-gas removal
process. It has, therefore, passed through a quench sys-
tem, which would remove the heavier oils, and the water-
gas shift reactor. In both cases, the water-gas shift
catalyst is also a hydrogenation catalyst so most of the
sulfur-containing compounds in the remaining oils should
be hydrodesulfurized. An analysis of the data presented
with the Hygas process indicates that the BTX stream will
contain perhaps 50 percent very light oil (similar to kero-
sene or No. 1 fuel oil), a small quantity of sulfur indicated
as H2S, and BTX. This by-product stream could be sold
111-26
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to a cooperating petroleum refinery that would mix it
with similar streams, and fractionate or solvent extract
it for BTX recovery, thus completing the desulfurization
process.
The tars and oils are differentiated only boiling
point. In the Lurgi process, the tars are separted from
the oils by condensation at an unspecified temperature.
Of the processes described in this report, only the Synthane
and Lurgi processes produce tars as by-products; however,
the tars and oils produced as by-products from the gasifica-
tion processes have characteristics similar to the liquid
product stream from the COED process. They may also
be similarly treated for recovery of their energy values
(in an environmentally acceptable manner).
The primary contaminant in the tars and oils is sul-
fur. In the tars, the sulfur may exist in the form of
organic sulfides, mercaptans, or thiophenes. The oils
might contain similar compounds but with lower molecular
weights, and, additionally, the oils might contain dis-
solved hydrogen sulfide and carbonyl sulfide. The inherent
characteristics of the coal feed, and the techniques of
primary treatment, will determine the total quantity of
tars and oils, the relative proportion of tars and oils, and
the quantity of sulfur contained in these organic liquid by-
product streams. For example, the Lurgi process,
operating on Navajo coal, produces a relatively large
quantity of tars, but these tars contain sufficiently low sul-
fur to be burned directly for the recovery of their heating
value. The relative quantity of sulfur in the tars is fur-
ther illustrated by the two examples presented for the
Synthane process. With an Eastern coal, the sulfur con-
tent of the tars is quite high (about 5 percent), but it is
lower with a mid-continent coal.
The oils may be quenched from the gas stream
either before or after the water-gas shift reaction stage.
Those oils which are condensed after contact with the
shift catalyst may be partially hydrodesulfurized because
of the hydrogenation characteristics of this catalyst, but
oils condensed before the shift reaction will also contain
sulfides, mercaptans, and thiophenes.
111-27
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Sulfur may be removed from both the tars and oils
by hydrodesulfurization. These organic liquids exist in
the raw gas stream because of coal pyrolysis and, there-
fore, they should be amenable to distillate hydrodesul-
furization techniques as discussed more fully in that pro-
cess synopsis. However, the tars and oils are the first
materials quenched from the raw gas and may, therefore,
contain particulates; residual HDS techniques, consequently,
may be more applicable. After this treatment, the tars
might be satisfactorily consumed by an electric utility
which also burns No. 6 fuel oil. The oil fraction is more
similar to No. 2 fuel oil which is generally consumed in
homes and industry. It can be assumed that this hydro-
desulfurization would not be performed onsite but, rather,
the raw tars and oils would be marketed by a cooperating
petroleum refinery for further upgrading.
The tars might be reinjected into the gasifier. In
some overseas installations of the Lurgi gasifier, coal
fines which are too small to be accepted by the fixed-bed
unit are briquetted with the tars so that they may be
gasified.* Similarly, the tars may be reintroduced into
the gasifier by spraying in the top zones above the coal.
In both these instances, the tars are eventually hydro-
cracked into lower molecular weight oils that are more
marketable.
The tars may be gasified by partial oxidation,
similar to the Texaco process. A low-Btu gas could be
made by air oxidation to fulfill a portion of the process
energy and steam requirements. It is understood El Paso
considered this alternative for its Lurgi system but re-
jected it in favor of air-gasification of coal. The tars
could also be oxygen-treated to produce a synthesis gas
for reintroduction into the primary product flow stream.
Briquetting is a relatively expensive operation, about $1/million
kcal (25£/million Btu), and is not planned for the Lurgi El Paso
project.
111-28
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Although the tars might be partially oxidized in a
separate processing step, it should be possible to rein-
troduce the tars directly into the primary gasifier in the
high-temperature zone, thus gasifying them directly.
Gasifier modifications and operating techniques, required
for this direct gasification, have not yet been developed.
(5) Gases—General Introduction
The primary gaseous pollutants generated in clean-fuel
conversion processes are:
• Sulfurous compounds
Hydrogen sulfide (H^S)
Carbonyl sulfide (COS)
Other reduced sulfurous compounds: for
example, mercaptans (RSH), sulfides (RS),
carbon disulfide (C$2)
Sulfur dioxide (SO2)
Elemental sulfur (82, 84, etc. vapors).
Carbon Oxides
Carbon dioxide (CO2), not normally considered
a pollutant
Carbon monoxide (CO).
* Light hydrocarbons
Methane (CH4)
Ethane (C2H6)
Ethylene (C2H4)
111-29
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Traces of higher aliphatics and olefins
BTX (benzene, toluene, and xylene vapors).
Nitrogeneous compounds
Ammonia (NHg)
Hydrogen cyanide (H CN)
Thiocyanates (SCN -)
These gases exist as by-products in some processes (Gas
Combustion Retort, Solvent-Refined Coal, HDS, and so forth),
but they may be part of the main process flow in others. If they
are by-product streams, the sulfurous and nitrogeneous com-
pounds must be removed so that the by-product may be used as
a fuel gas. When the gaseous stream is the primary product,
the sulfurous and nitrogeneous compounds should be removed
for pollution control purposes while achieving maximum re-
covery of hydrocarbons. If a high-Btu product is to be manu-
factured, the carbon monoxide and hydrogen will be converted
to methane, but water vapor and carbon dioxide must be rejected
from the system with minimum contamination by pollutants.
In general, therefore, the aims of gas purification can be
listed as follows:
Rejection of sulfurous compounds from the system,
primarily as solid elemental sulfur with minimum
evolution of gaseous sulfurous compounds.
Rejection of nitrogeneous compounds from the
system, preferably as by-product ammonia.
If the gaseous stream is to be used directly for fuel
value, water vapor and carbon dioxide need not be
removed from the system. This will minimize
hydrocarbon losses.
If the gaseous stream is to be converted to hydro-
gen (for example, for HDS or hydrocracking feed
gas), the carbon monoxide must be shifted to car-
bon dioxide. This carbon dioxide, in addition to
111-30
-------
the hydrocarbons present, must eventually be
removed from the system.
• If the product gas is to be used as a synthesis gas
(for example, for manufacture of ammonia, methane,
alcohol, or oils by a Fischer-Tropsch reaction),
carbon dioxide and water vapor must be eliminated
from the system and sulfur must be removed to
very low levels to protect the sensitive downstream
catalysts.
The most complicated purification steps are employed in
the high-Btu gas manufacturing processes using coal. These
are tabulated below:
• The majority of the water and oil vapors are re-
moved by cooling and condensation.
Water-washing is used to remove water soluble
contaminants such as ammonia, cyanides, and
phenols.
• Oil-washing is used to recover BTX and oil vapor.
• Special chemical sorption techniques are used to
remove acid-gases such as I^S, COS, HCN, and
C02.
Adsorption on special beds is used to eliminate
trace quantities of pollutants.
The four processing steps that require the greatest
attention for pollutant control are water-washing, acid-gas
removal, sulfur recovery, and trace pollutant removal. These
are discussed in detail below.
1. Water -Washing
Water-washing is use<| to remove ammpnia, phenolic
compounds, and cyanide fr^fh the primary gas stream. The
treatment of foul water from the gas-scrubbing operation
is not well defined in the processes studied in this report.
111-31
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Only the HDS and Texaco Partial-Oxidation processes
have been developed to the point where this problem has
been analyzed; however, both of these processes operate
on petroleum feedstocks with low nitrogen content and,
therefore, yield little ammonia or cyanide. The Lurgi
process, as designed for the El Paso Natural Gas Com-
pany, has been engineered for wastewater treatment and
disposition of potential pollutants which would be conden-
sed and washed out of the process gas. It has not yet
been installed, however, so the operation of this section
of the plant cannot be empirically verified. The flow
sheet presented for the Hygas process indicates the
techniques that would be used for wastewater treatment,
but this system is not yet completely engineered. The re-
maining processes analyzed have not defined their waste-
water treatment. This problem might well be an appro-
priate subject for an in-depth study, with particular
emphasis on treating wash-liquors drawing on experience
gained from dealing with coke-oven gases.
The phenols which are evolved in the primary
reactor are water-washed from the process gas stream.
They are recovered directly from wash water (as was
discussed in an earlier section on wastewater treatment).
Ammonia was manufactured in the primary reaction
stage of the process by the action of hydrogen upon the
nitrogen in the raw material. Some oil feedstocks have
essentially no nitrogen content, and ammonia will not be
present in the product gas. On the other hand, the kerogen
in shale is very high in nitrogen and will make considerable
quantities of ammonia. Coal-gased systems will con-
tain intermediate quantities of nitrogen and, there-
fore, similar quantities of ammonia in the product
gas.
Hydrogen cyanide probably is not formed directly
in these processes but rather is the product of a secondary
reaction of ammonia and hot carbon. The quantity of
this cyanide in the gaseous stream will be a function of
the amount of ammonia that is present and in contact with
the feed at temperatures hot enough to form cyanide.
111-32
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In the Hygas and Lurgi processes illustrated, the
water-wash follows the shift reactor. The catalyst in this
reactor hydrogenates a large portion of the cyanide into
ammonia. If the ep+ire process stream passes through the
shift catalyst, the cyanide treatment problem would be
minimized because nearly all the cyanide would be converted;
however, both these illustrated processes use a split-shift
configuration. In this arrangement, part of the raw gas is
bypassed around the shift reactor and the remainder of the
gas is shifted. The resultant combination stream there-
fore contains some HCN that is removed in water wash-
ing.
As was discussed earlier, under liquid waste disposal,
the cyanides can be converted to thiocyanates in the presence
of sulfur in an oxidizing atmosphere. This conversion is not
expected to take place within the primary gas stream, but it
may take place in subsidiary streams in the plant.
The third major pollutant that may be removed by
water-washing is phenol. In general, phenol and higher
hydroxylated aromatics such as cresols, are formed
directly during the devolatization of coal. Therefore, their
occurrence in the gas stream is a function of the character-
ization of the coal feedstock utilized and the initial pro-
cessing conditions employed.
The water wash will also dissolve some hydrogen
sulfide, carbon dioxide, and carbonyl sulfide from the
product gas as well as small quantities of hydrogen, car-
bon monoxide, and methane. Most of this material will
be recovered from the water when it is flashed to atmos-
pheric pressure. Additional sulfur is recovered when the
ammonia liquor is steam stripped. In the Lurgi process
an ammonia liquor with 20 percent NH3 is recovered from
the steam stripper and sold. Cyanides pass, with the acid-
gas, to the sulfur recovery unit where the cyanides are
converted into thiocyanates for disposal.
The Hygas process assumes complete stripping of the
wastewater similar to a system used by United States
Steel engineers in its treatment of coke-oven gas, leav-
ing minute quantities of H2S, HCN, and NH3 in the water.
111-33
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The overhead gases from the stripper are treated for
recovery of 'anhydrous ammonia and then fed to an incin-
erator prior to arrival at the final sulfur recovery unit.
Alternatively, the ammonia-free, scavenged gas could be
recompressed and returned to the primary process gas
stream, possibly before shift conversion to eliminate
HCN from the system.
2. Acid-Gas Treatment
After the raw gas has been water-scrubbed to remove
water soluble compounds such as ammonia, cyanide, and
phenols, and perhaps oil-scrubbed to remove oily vapors
from the gas, it undergoes acid-gas scrubbing to remove
sulfurous compounds and carbon dioxide from the gas
stream. In the case of gasifying coal to produce a high-
Btu gas, the sulfur content must be reduced to very low
levels to minimize poisoning of sensitive downstream
methanation catalysts. Carbon dioxide must also be re-
moved because it would dilute the product gas and compete
with the carbon monoxide, in the methanation reaction,
resulting in higher hydrogen consumption.
Acid-gas scrubbing may be employed before the
water-gas shift reaction, if conventional, sulfur-sensitive
shift catalysts are employed. Overall process efficiency
may be improved if the newer, sulfur-tolerant shift
catalysts are used because cooling and condensation, re-
quired before acid-gas treatment, is postponed until after
the shift reaction and makeup high-pressure steam to the
shift reactor is minimized. Carbon dioxide is also manu-
factured during the shift reaction. With this delayed acid-
gas treatment, (after shift conversion), only a single
stage of CC>2 removal is required.
The carbon dioxide, which is to be removed in the
acid-gas treatment, is manufactured at several points of
the process. In many processes, the endothermic heat
of the steam-coal reaction is supplied by burning a portion
of that coal thereby yielding some carbon dioxide (for
example^ in the steam-oxygen gasifier of the Hygas process).
Some carbon dioxide is also generated in the primary
111-34
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gasification stages and still other carbon dioxide is
generated in the water-gas shift reactor by conversion of
some of the carbon monoxide in the process gas stream.
The sulfurous compounds, to be removed during acid-
gas treatment, are generated during the gasification of the
coal. The primary purpose of most of these processes is
to cbnvert a high-sulfur fuel, such as coal, into a low-sulfur
energy source, such as substitute natural gas. The sulfur
which occurs naturally in the feedstock must be eliminated
from the system.
The primary sulfur compound which exists in the pro-
cess gas stream is hydrogen sulfide (E^S). When carbon
monoxide is present in the process gas stream, carbonyl
sulfide (COS) will also be generated. If a portion of the
process exists at relatively high temperatures, carbon
disulfide (CS2) will be formed. If the hydrogen and carbon
monoxide concentrations in the primary reaction stage are
approximately equal, the thermodynamically expected
carbonyl sulfide concentration is about 1/30 of the hydrogen
sulfide concentration, and the carbon disulfide concentration
will be even lower. As will be shown later, however, hydro-
gen sulfide is more readily removed from the process gas
stream than COS or CS2« Therefore, the processing con-
ditions which affect the manufacture and conversion of COS
and CSp must be examined to determine the overall pollutant
discharge from the facility.
Other forms of sulfur such as organic sulfides, mer-
captans, and thiophenes, are also produced in the primary
reaction step during the devolitization of coal. These
heavier sulfur compounds will be collected with the tars
and oils from the process and then be hydrodesulfurized
as discussed earlier. The I^S, COS, and CS2 are the
primary sulfur types to be found in the process gas as it
Breaches the acid-gas treatment section of the facility.
One of the major problems in all of the clean-fuel
processes is accurate quantitative measurement of the
sulfur. Sulfur compounds, particularly H2S, tend to
sorb onto metal surfaces, reducing the accuracy of the
sulfur analyses. Therefore, the data presented in the
111-35
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process synopses must be regarded as preliminary until
the processes are demonstrated on a reasonable scale.
The initial carbonyl sulfide concentration might be
one to five percent of the hydrogen sulfide concentration.
This depends upon the H2/CO ratio in the primary reac-
tion unit, the operating temperature of that unit, and the
kinetic approach to equilibrium in the H2S-COS system.
In the Lurgi, Hygas, and Texaco Partial Oxidation pro-
cesses, the acid-gas treatment unit is shown to follow the
shift-reaction stage. The shift catalyst used for these
sulfur-laden streams is active for the H2S-COS equilibrium;
COS + H2 -?H2S + CO.
The COS concentration should fall to about one to two per-
cent of the HoS concentration if the Ho/CO ratio is about
& £t
3:1. The Hygas and Lurgi flow sheets show this equilibra-
tion, but the Texaco flow sheet does not. If this catalyst
is indeed active for the HgS-COS system, the use of delayed
acid-gas treatment (after shift conversion) will enhance the
sulfur recovery from the process.
Two primary problems must be faced in the acid-
gas cleanup unit:
The sulfur must be removed from the process
gas stream to very low levels in order to pro-
tect the sulfur-sensitive catalysts.
The sulfur present in the regenerated off-gas
from the system must be present in the proper
concentration and form to permit economical
recovery as elemental sulfur.
These two constraints work at cross-purposes in most
acid-gas treatment systems; complete sulfur removal from
the process gas requires high carbon dioxide removal.
The sulfur concentration in the process gas is only about
four percent of the carbon dioxide concentration after shift
-conversion; therefore, the resulting regenerated gas will
be too lean in sulfur for economical operation of a conven-
tional Claus system (discussed later) for elemental sulfur
recovery. In these cases, the sulfur must be recovered
by a technique similar to the Stretford process but the
Stretford process does not remove COS or CS2 from the
gas (as discussed later in this section).
II1-3 6
-------
One approach to the problem of insufficient hydrogen
sulfide concentration is the natural tendency of some acid-
gas removal processes to absorb H^S selectively over
CO2. This natural selectivity can be employed in a
double-sorption system to recover a high-concentration
H2S stream for a Glaus plant feed and a CO2-rich stream
for disposal. However, this approach also has short-
comings as will be discussed.
The following paragraphs discuss the alternative
types of sulfur removal processes in general terms. The
subject of elemental sulfur recovery via the Claus system
is then presented, followed by a discussion of Claus off-gas
purification for emission control and sulfur guards for
removing traces of sulfur that would otherwise poison
the sensitive downstream catalysts.
(6) Gases—Sulfur Removal Processes
Processes for removal of sulfur from acid-gases may be
divided into four general categories according to their mode of
operation:
Physical solvents; some liquids have natural sol-
vation capabilities for hydrogen sulfide and carbon
dioxide
Alkaline salt solutions; alkaline solutions will
react chemically with the acid gases, permitting
removal
Amines; amines will react with acid-gases,i
permitting removal
Indirect oxidation; some combinations of chemicals
will absorb hydrogen sulfide and recover it directly
as elemental sulfur by oxidation in the regeneration
step.
Several other approaches have been suggested for acid-gas
removal, but are not commercially applied.
111-37
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1. Physical Solvent Processes
Physical absorption is the dissolution of an acid-gas
in a solvent without undergoing any chemical transforma-
tion. Hydrogen sulfide and carbon dioxide are highly
soluble in many organic solvents. These solvents can
also remove COS, CS2, and mercaptans; but unfortunately,
.they also remove some higher hydrocarbons, particularly
aromatics, as well. Therefore, the regenerated gases
which are produced by flashing these solvents to lower
pressures may not be suitable for Claus-type sulfur dis-
posal facilities, unless the solvent off-gases can be
fractionated to recover these aromatics.
.Some of the well known solvent-based processes
are defined in Table III-8. In these processes, acid-gas
is absorbed at elevated pressure and ambient temperature
by countercurrent scrubbing with solvent. (The Rectisol
process, however, operates at sub-ambient temperatures
using refrigeration). The rich solvent is flashed at an
intermediate pressure and the evolved gas, chiefly methane,
is recompressed and returned to the absorber. Additional
lower pressure stages of flashing may be used before the
final stripping.
The solvent-based processes find their greatest
application when system pressure is high. The solubility
of the gas is approximately proportional to pressure, and
solvent-based processes become economically attractive
at high partial pressures of the acid-gases.
The solubility of hydrogen sulfide and carbon diox-
ide are not the same in the solvents shown in Table III-8.
Usually hydrogen sulfide is significantly more soluble.
Therefore, these processes often offer the option of con-
centrating the H2S stream as a potential feed to a Claus
.unit. This may be accomplished by two-stage absorption
where the first stage absorbs most of the H^S with re-
duced CC>2 removal, or selective stripping may be used
with a single sorption stage. The advantages of the
solvent-based processes are:
High solubility of solvent for H S and CO
£t £t
111-38
-------
vo
Table III-8
Solvent-Based Process
V
Process
Purisol
Fluor
Selexol
Rectisol
Sulfinol
Developer
Lurgi
Fluor
Allied
Lurgi
Shell
Solvent
N-Methyl-2
pyrrolidone
Propylene carbonate
Dimethyl ether
polyethylene glycol
Methanol
Tetrahydrothiopene
1-1 dioxide (Sulfolene)
plus diisopropanol-
amine (DIPA)
Typical Treated
Gas Purity
H2S CO2
4 ppm 2-3%
Above 1-2%
4 ppm
4 ppm 2-4%
3 ppm 60 ppm
4 ppm Below
0.5%
Attainable Treated
Gas Purity
H2S
2 ppm
1 ppm
1 ppm
CO2
10 ppm
0.5%
10 ppm
200 ppm
-------
Minimum consumption of heat, cooling, and
power
Minimum corrosion because of low tempera-
tures
Simultaneous removel of trace impurities
such as HCN, COS, and organic sulfur com-
pounds
Excellent thermal and chemical stability of
the solvents
Possible selective recovery of E^S.
Disadvantages of the solvent-based processes include:
The solvent is costly, and losses must be
minimized
The impurity must have high partial pressure
Thorough impurity removal is difficult
High methane and ethane solubility adds
recompression costs or potential loss of
product
Heavy oils (and water for some solvents) are
readily absorbed, requiring changes to the
process flow scheme or preremoval before
treating.
2.. Alkaline Salt Solution Processes
The use of alkaline salts, such as sodium or potas-
sium carbonate, for the absorption of H2S and CO2» dates
back to a German patent in 1804. The Seaboard and
Vacuum Carbonate processes were developed, but the
true potential of the carbonates was not realized until
the Bureau of Mines performed an extensive study about
20 years ago. The hot potassium carbonate, or "hot pot"
IH-40
-------
process, is primarily utilized for the removal of carbon
dioxide; H^S is not removed from acid-gas unless CC>2
is also present. Hot carbonate processes are also re-
stricted to high absorber pressures and high acid-gas
content streams because the gases must physically absorb
into the solution before they can chemically react. The
CO2 content of the treated gas stream may be as low as
0.1 percent, although 2 percent residual CC>2 content is
usually practiced due to economic considerations. The
residual t^S content of the process gas may be as low as
10 ppm.
In this process, CC>2 and H^S combine with potas-
sium carbonate at temperatures of 80 to 120 degrees C
(180 to 250 degrees F). Regeneration is accomplished by
steam stripping the solution at the absorption temperature
but at lower pressure. The small temperature difference
between the absorber and the stripper results in lower
utility requirements than many other processes. Capital
costs are also reduced.
Proprietary processes, such as the Benfield,
Catacarb, and Giammarco-Vetrocoke (CC>2), are based
on catalytic improvements of the hot carbonate process.
The catalyst increases the rate of absorption and stripping
of solution, resulting in lower solution flow rates, lower
heat requirements with better acid-gas removal, and a
consequent reduction in capital costs for plant.
The hot carbonate processes offer one distinct ad-
vantage over the others discussed here: carbonyl sulfide
and carbon disulfide are hydrolyzed during treatment;
COS + H0O £ H S + CO •
Ct £t £t
Therefore, these compounds are converted to hydrogen
sulfide which is absorbed by the solution. Only H^S,
therefore, is regenerated with the carbon dioxide. In
the synopsis of the Texaco Partial Oxidation process, a
,"hot-pot" process was arbitrarily included in the flow
sheet variation for syn-gas production, to show the effect
of COS hydrolysis.
Another benefit of the "hot-pot" process is its natural
partial selectivity. Benfield, a process developer, will de-
sign a two-stage treatment system which generates an H S-rich
III41
-------
gas which is a suitable feed for a Glaus plant and a
gas claimed to contain only 50 ppm IH^S. The primary pro
cess problems are:
Corrosion
Erosion
Foaming
Solution degradation by formate* and sulphate
• • formation
3. Amines
The best known regenerable acid-gas removal
processes employ amines for the complexing of H^S or
CC>2 from the gas stream. These processes, introduced
about 40 years ago, have been criticized because of their
high heat requirements, corrosion problems, and foam-
ing tendencies. Every new process introduced in recent
years has been compared favorably to amines; ;yet amine
processing continues to be the most common type of new
acid-gas installation. Amine processing must, therefore,
still be considered as a viable acid-gas removal system.
Some of the active amine processes are defined in
Table III-9.
In a conventional amine process, a 15 to 25 percent
solution of monoethanolamine (MEA), or diethenolamine
(DEA) in a water or glycol base is used to contact the
acid-gas. The acid constituents are removed in a
compound-complex formation, and the resulting solution
is stripped of these constituents by heating. Advantages
of the MEA process are high reactivity, low solvent cost,
ease of reclamation, good stability, and low hydrocarbon
content in the acid-gas produced. The plant investment
is often lower than for other processes. However, the
MEA process cannot satisfactorily process streams with
A salt or ester of formic acid (CH O ).
£ £i
111-42
-------
Table III-9
Amine Processes
OJ
Process
Amine
Economine
ADIP
Active Component
Monoethanolamine
(MEA)
Diethanolamine
(DBA)
Diglycolamine
(DGA)
Diisopropanolamine *
(DIPA)
Solution
15% MEA in
water
25% DBA in
water
50-70% DGA
in water
Typical Treated
Gas Purity
H2S
Below
4ppm
Below
400 ppm
4ppm
CO2
100 ppm
100 ppm
Attainable Treated
Gas Purity
H2S
1 ppm
50 ppm
4 ppm
CO2
20 ppm
200 ppm
100 ppm
*Often used in conjunction with Sulfinol solvent
-------
COS and CS 2 because of degradation of the MEA. It is
also ineffective in removing mercaptans and has higher
vaporization losses than other amine processes.
The DEA process has lower reactivity and higher
solution cost than the MEA process but is resistant to
degradation from COS and CS2. Higher amines offer
partial selectivity for H^S over CO2, but they have not
generally been employed. More recently, diisopropanol
amine (DIPA) has been offered which attains higher
selectivity. Diglycolamine (DGA) is also available and
may be operated at higher sorbent concentrations and,
therefore, lower investment and utility costs.
In the process synopses presented in Section: 2 of
the report, MEA processing is used only in the purifica-
tion of gases resulting from hydrodesulfurization of crude
oils. In this operation, the oils are treated with pure
hydrogen; therefore, the only sulfur compound in the off-
gas is hydrogen sulfide for which MEA processing works
satisfactorily.
Although DEA was originally specified as the acid-
gas cleanup process for the Hygas process pilot plant,
this system has now been converted to DGA. However,
the Hygas process synopsis presented in Section II
utilizes a solvent-based acid-gas cleanup system (dis-
cussed earlier).
Tertiary amines (TEA and MDEA) were used at the
old Bureau of Mines' demonstration plant for coal gasifi-
cation and liquefaction. They are still specified for puri-
fication for Koppers-Totzek process effluents but are not,
however, used significantly in commercial installations.
Diisopropanol amine is used extensively in con-
junction with the Sulfinol solvent process. The combina-
tion of solvent and amines apparently yields excellent
acid-gas removal with a reasonable selectivity of H2S
over CO2.
The amines remove some COS from the gas; how-
ever, their action is not a strict acid-gas reaction. For
IIM4
-------
some light amines, the COS-amine compound is stable,,
but with heavier amines, the amine can be reclaimed by
hydrolysis at elevated temperatures, regenerating the
COS as H2S.
4. Indirect Oxidation Processes
A series of processes has been developed in recent
years for acid-gas purification with elemental sulfur as
the final product. In the indirect oxidation schemes, the
hydrogen sulfide in the gas reacts selectively with some
constituents in the scrubbing liquor and is removed from
the gas stream. During regeneration, air is blown through
the solution and sulfur oxidizes to the elemental form. It
is collected as a froth on the top of the solution and re-
covered by filtration or centrifugation. Many of these pro-
cesses have had limited commercial use in this country.
Two processes that deserve consideration are shown in
Table III-10. Process advantages claimed for these sys-
tems are that they:
Treat the sulfur selectively
Can usually remove the sulfur to low levels
in the treated gas
• Produce an elemental sulfur product, .elimi-
nating the need for a Claus unit.
Table III-10
Indirect Oxidation Processes
Process
Giammarco-
Vetrccoke
Stretford
Solution
KsAsOs ;
Na2CO3 +
Na2VO3 +
organic oxida-
tion catalyst
Treated Gas Purity
ti2s
1 ppm
1 ppm
CO2
May be removed
No change
IIW5
-------
They generally require a large reactor, however, because
of low rates of sulfur reduction in the regenerator unit.
The use of arsenic-containing solutions has, in some in-
stances,posed drawbacks to the acceptance of some pro-
cesses; insufficient removal of arsenic from the by-
product sulfur renders it unfit for sale.
Another potential problem is the removal of COS
and CS2. Processes like Stretford are selective for H2S
removal. Thiosulfate formation is also possible and
bleed streams would contain some high-cost vanadium
catalysts.
(7) Gases—Elemental Sulfur Removal from B^S Streams
The processes available to recover elemental sulfur from
H2S-rich gas streams and the treatment necessary to clean any
tail-gases which may be present is discussed below.
1. The Glaus Process
The conventional technique for disposal of the hydro-
gen sulfide stream from the acid-gas treatment is the
Glaus process. In this scheme, the H^S is reacted with
controlled amounts of oxygen over bauxite catalysts to
form elemental sulfur. In some variations, one-third of
the feed stream is burned to sulfur dioxide, and the re-
maining stream of E^S is mixed with the SC>2 to provide
the proper catalyst feed. In well-designed, multistage
Claus units, sulfur removal efficiencies of 95 percent
can be attained if the feed gas contains high-H2S concen-
trations. Efficiency falls as the hydrogen sulfide concen-
tration in the feed gas is decreased; for example, 80 to
89 percent efficiency is achievable at 10 percent H2S
concentrations in the feed, and 70 to 85 percent efficiency
raay.be realized with 5 percent H2S streams. Higher
efficiencies are limited by the thermodynamics of the
process. The sulfur is formed in the vapor state and
the reversible process is governed by equilibrium.
IIM6
-------
The Glaus unit feed must contain enough sulfur so
that the heat of reaction is sufficient to sustain the flame.
If the sulfur concentration is low, the problem is similar
to trying to burn a f^sl mixture which is too lean. Even
direct addition of oxygen can be insufficient with an E^S-
lean feed because excessive amounts of CC>2 may quench
the flame. This requirement for high-HgS concentration
in the Claus unit feed is the reason sulfur-selective puri-
fication processes have been stressed in the preceding
paragraphs.
If possible, hydrocarbons should be eliminated from
the feed to the Claus unit. Although these are burned to
provide heat (and are occasionally added intentionally if
the feed sulfur concentration is low), they also react with
the sulfur in the gas to produce COS and CS2. These com-
pounds are not oxidized in the conventional Claus plant,
although special, high-temperature stages may be added to
the system to minimize COS losses. For these reasons,
Claus plants are not generally satisfactory following many
solvent-based acid-gas removal processes. The solvents
also remove some hydrocarbons from the gas stream
resulting in the same problems.
As was discussed in earlier paragraphs, hydrogen
sulfide may be oxidized directly to elemental sulfur in the
Stretford process. This system will operate on reduced
concentrations of H2S, below the range of a Claus system;
however, the Stretford system cannot eliminate COS and
CS2 from the waste gas. Recent undisclosed process ad-
vances claimed by the licensor of the Stretford process
should significantly reduce the cost of this unit process,
and it may now be competitive with the total of a Claus
plant with its necessary tail-gas purification unit.
2. Claus Tail-Gas Purification
The sulfur content in the effluent from a Claus unit
may be relatively high. If the Claus tail-gas is incinerated,
sulfur will exist in the stack gas as sulfur dioxide. Sulfur
dioxide removal is now considered to be a viable process
for power plant stack gases using limestone scrubbing.
111-47
-------
However, other processes, discussed below, may be more
desirable for purification of Glaus plant tail gas.
The Wellman-Lord system has been satisfactorily
applied to Claus plant tail-gas. In this scheme* the Claus
effluent is incinerated, converting the reduced sulfides and
elemental sulfur in the stream to SC>2. This sulfur dioxide
is concentrated in a sodium sulfite-bisulfite system and
regenerated, in a concentrated form, for feed back into the
Claus unit.
At least three systems have been developed that
operate on hydrogen sulfide, (for example, Beavon and
SCOT processes). The total sulfur content of the Claus
effluent is catalytically hydrogenated, using reducing gas,
to hydrogen sulfide. * In two process variations, the H^S
is treated by a Stretford process while a third variation
process uses a selective amine for I^S recovery recycling
it as Claus feed.
The licensors of all four of these Claus tail-gas
purification processes claim that the effluent will be equiva-
lent to less than 250 ppm SO2. However, three of the
processes might require incineration of the tail-gas before
venting it to the atmosphere.
Several other processes have been developed to treat
Claus off-gases. These other processes do not yet have the
commercial experience of the four discussed above.
(8) Gases—Trace Sulfur Removal from Major Streams
Final purification processes to remove the last traces of
sulfur from the main stream are discussed below.
1. Sulfur Guards
Trace quantities of sulfur impurities may be removed
from the process gas stream by absorption of the surface
or in the pores of some solid materials. These adsorption
processes are usually performed with fixed beds of adsorbent.
* Apparently, the operating conditions do not promote the water-
gas shift reaction, the hydrogeneration of
HI-48
-------
When the adsorbent capacity to retain the impurity is
reached, the adsorbent may be discarded or, more com-
monly, regenerated by heating. Usually, two beds of
solids are used when regeneration is employed: one being
used for adsorption while the other is being regenerated.
Sulfur guards will be required in all processes pro-
ducing high Btu synthetic gas to protect the sensitive
downstream catalysts. Even a few ppm of sulfur may
poison these materials. For conventional methanation
processes using nickel catalyst, catalyst lifetime will be
about two years when the gas contains 0. 07 ppm sulfur;
but a catalyst life of five years or more is expected with
sulfur-free gas.
2. Activated Carbon
To attain very low-sulfur concentrations, the pro-
cess gas can be passed through a fixed bed of activated
carbon which has been impregnated with certain metal
oxides. Hydrogen sulfide as well as carbonyl sulfide,
carbon sulfide, thiophenes, and mercaptans are removed
by the carbon. When the capacity of the bed is reached,
it is removed from service and regenerated with steam
and air. Some elemental sulfur formed during regenera-
tion accumulates in the carbon pores and eventually de-
activates the bed. Therefore, after repeated adsorption-
regeneration cycles, the impregnated carbon must be
replaced with fresh sorbent.
3. Iron-Sponge Treatement
Iron-sponge treating may be the most economical
sulfur recovery system if the process gas contains less
than 300 ppm of sulfur and CC>2 removal is not necessary.
Iron sponge is prepared by impregnating wood shavings
with hydrated iron oxide. The mixture contains about
112 kg of iron oxide per cubic meter (7 Ib/ft^). Sponge
is loaded into the treating chamber and process gas con-
taining traces of impurities is fed at ambient tempera-
tures. The bed moisture content is held at 30 to 50 per-
cent and the pH is alkaline. At 100 percent efficiency,
III49
-------
0. 64 kg (or Ib) of H2S will be removed per kg (or Ib) of
iron oxide. Spent iron oxide is often replaced with fresh
material, although it may be regenerated by air oxida-
tion.
The advantages of iron sponge treatment are that it:
Completely removes even low concentrations
of H2S
Is effective at any pressure
Requires small investment
The disadvantages include the fact that it:
Requires frequent change of spent sponge
Is a batch process
4. Zinc Oxide Treatment
Zinc oxide has a special affinity for sulfur; the
equilibrium partial pressure of hydrogen sulfide over
zinc oxide is measured in parts per billion, even at
elevated temperatures. Zinc oxide beds have been used
for many years as the final step for trace sulfur removal
from gas streams.
5. Caustic Washing
Washing the gas with caustic will remove sulfur
compounds to extremely low levels. However, carbon
dioxide is also removed in this process, so caustic wash-
ing is generally applied only when the CO2 concentration
is also quite low. This system is often used for final
trace sulfur removal if there is a ready means of dis-
posal of the spent caustic.
HI-50
-------
(9) Gases—Sulfur Dioxide Removal from Combustion Gases
The primary source of sulfur dioxide emissions in clean-
fuel processes is from combustion. This combustion might
occur in three major areas of the process:
Power production—Raw or gasified and desulfurized
fuel may be used for generating steam and electric
power required for most of the processes considered.
This fuel will contain some sulfur which will appear
in the stacks as sulfur dioxide.
Incineration—In some processes, sulfur-containing
streams are incinerated to destroy combustible
matter and convert sulfur to the more innocuous
form of sulfur dioxide.
Calcination— In the CO^ Acceptor process, calcium
carbonate is burned with gasifier char to regenerate
the activated lime needed in the process. The char
will contain small quantities of sulfur which will
report to the stack gas as sulfur dioxide.
The calcination problem, in this study, is unique to the CC>2-
Acceptor process and is discussed in detail in that process
synopsis. General discussions of the other two sources of
sulfur dioxide are presented below.
Incineration is specifically encountered in two of the
process synopses. In the Lurgi system, incineration is speci-
fied for the off-gas from the Stretford unit which operated
on the H2S-rich stream of the acid-gas cleanup system. The
discharge of the Claus plant, in the Hygas process, was incin-
erated along with a scavenged gas, prior to being fed to a
Wellman-Lord sulfur recovery unit. Note that none of the
process developers suggest incineration of the bulk CC>2 stream
which may exist in the process.
The incineration practiced in the Lurgi process merely
reduces the sulfur compounds and converts them to sulfur
dioxide. The quantity of sulfur dioxide contained in this stream
is small, well below an acceptable concentration for venting
to the atmosphere, based upon permissible concentration of
SO in stack gases.
111-51
-------
The sulfur concentration in the off-gas from a Claus plant
is relatively high. Incineration is practiced, in this instance,
if a Wellman-Lord system is used for the tail-gas cleanup proc-
ess. As was discussed earlier, the Wellman-Lord system
operates on sulfur dioxide, concentrating it for return to the
Claus unit. The sulfur dioxide content of the effluent from the
Wellman-Lord system is sufficiently low to be discharged to the
atmosphere.
The primary source of sulfur dioxide emissions in these
processes is the stack-gas manufactured during the production
of steam and power for the facility. In some processes, such
as Synthane, a gasifier char is burned directly for steam genera-
tion. The sulfur content of this char is relatively high and the
stack gases from the power plant would contain an unacceptably
high sulfur dioxide concentration. Consequently, stack-gas
cleanup processes were applied in these cases to reduce the
SC>2 emissions from the boilers to less than 2.16 kg SC>2/
106kcal of the feed fired (1.2 Ib SO2/106 Btu), the present
EPA New Source Performance Standard. It will be noted, in
the process synopses, that stack gases are still the major source
of sulfur emissions from the facility, even after application of
approved cleanup processes to meet the specified effluent levels.
An alternative approach for power generation is the gasi-
fication of the char fuel and desulfurization of the fuel gas before
combustion. This approach was specified in the Lur.gi and in the
Hygas process synopses. It was also applied in this study as an
alternative means of power generation for the Synthane process
and others. This technique reduces the sulfur dioxide emissions
from this part of the plant by a factor of 10 and also increases
the overall sulfur recovery from the facility. Gasification of the
raw fuel and desulfurization of the fuel gas is the basic approach
used for low-Btu gasification and combined-cycle power genera-
tion. It is expected to have significant application in future
electrical power plants.
TII-52
-------
3. INTEGRATION OF SULFUR REMOVAL PROCESSES INTO THE
CLEAN FUEL FLOW SCHEMES
The application of the alternative pollution control techniques
discussed in the previous section to the clean fuels processes analyzed
in this study varies from process to process. Some of the process
developers have not included pollution control systems in their pro-
cess designs, and it was necessary to apply an acceptable alternative
in these cases. But it was not possible within the scope of the study
to ensure that the choices made were optimum, from an economic
and pollution control standpoint, for the waste streams considered.
When process developers had specified a pollution control system in
their design, the analysis of costs and material balances was made
with that system.
The simplest sulfur- recovery system illustrated in the process
synopses presented in Section 2 of this report is that for the treat-
ment of the low-Btu gas produced by the Lurgi process. In this opera-
tion, the sulfur is removed directly from the gas by high-pressure
treatment with a Stretforci system. Carbon dioxide, carbonyl sulfide,
carbon disulfide, and other organic sulfur forms are not removed by
the Stretford process. The product gas, in this case, is consumed
directly and the residual SC>2 concentration in the stack is a factor of
10 lower than the present Federal permissible emission levels, although
this degree of control was required for the more stringent New Mexico
standards. If the intent in this process was to make a high-Btu gas,
the CC>2 must also be removed. The acid-gas treatment processes
which would remove the CO2. however, would also remove residual
sulfur compounds. This could result in a sulfur concentration in the
CC>2 vent gas of about 300 ppm. Even higher sulfur concentrations in
this waste gas would be expected with higher sulfur-content feedstock
or more efficient gasifiers that generate less
The traditional route for acid-gas cleanup is illustrated in the
HDS process. Here, only ^S is manufactured during the hydro-
genation of the sulfur compounds in the oil and this can be removed
by MEA scrubbing. The ^S concentration in the MEA regenerator
off -gas is high, so a Claus system can be employed satisfactorily.
After tail-gas purification, less than 250 ppm of sulfur compounds
would be exhausted to the atmosphere.
The Synthane process illustrates the application of nonselective
sulfur removal in the acid-gas system. In this case, a hot potassium
carbonate system was used for simultaneous HgS and CC>2 removal.
111-53
-------
The off-gas from the regenerator of this system contains only 1. 5 per-
cent H2S; therefore, a Stretford unit was used to recover this sulfur in
the elemental form. A Glaus unit would be highly inefficient given the
low concentration of H2S. In general, a Stretford may be expected to
allow less than 250 ppm of sulfur compounds in its effluent or the total
COS + CS2 in its feed, whichever is greater. Much better performance
might be expected in this instance because the only sulfur compound
expected in the off-gas from a "hot-pot" process is hydrogen sulfide.
Nevertheless, the process licensors' expectations of 99 percent re-
moval (150-250 ppm) have been quoted. The quantity of carbon dk>xide
in this exhaust stream is very high, approximately 7.1 million m /day
(250 million ft /day), so incineration would require expenditure of
excessive fuel.
Both the Texaco partial oxidation process (for hydrogen genera-
tion) and the Hygas hydrogasification process employ selective sulfur
recovery systems (two-stage, solvent-based systems). The operating
conditions of the first absorption stages are set to achieve maximum
sulfur removal. The sulfur concentration in the exhaust from this
unit, therefore, is high enough to feed a Glaus plant. In the second-
stage of acid-gas treatment, the remainder of the carbon dioxide is
removed with the residual sulfur. The H^S-rich stream passes
through a Glaus unit, followed by tail-gas purification, and the CO2-
rich gas stream is vented. Although the COp-rich gas stream may
contain very low H2S concentrations (less than 10 ppm), the relative
amount of carbonyl sulfide in these streams may be greater. As much
as 50 to 80 percent of the total carbonyl sulfide may report to this CO2
vent stream. Depending upon the sulfur level in the coal and the effi-
ciency of gasification, this stream may contain several hundred parts
per million of total sulfur compounds. Again, incineration would re-
quire significant quantities of energy because of the volume of this
gas stream (perhaps three to five percent of the product gas).
The sulfur removal system designed for the Lurgi process also
uses sulfur-selective solvent-based processes. The first stage of the
two-stage system produces an off-gas containing slightly less than
10 percent .sulfur. The total sulfur content of the COo-rich stream is
about 0. 75 percent. Both of these streams pass through separate
Stretford units for H2S recovery. The volume of the Stretford off-gas
from the original H2S-rich stream is low enough that it can be incin-
erated, but the volume of off-gas in the CO2-rich stream is too high
for incineration to be acceptable. Again, the Stretford process does
not remove carbonyl sulfide; a COS concentration of about 80 ppm is
exhausted with this vent stream. Higher COS concentrations would
be expected if the initial coal feedstock contained more sulfur.
Ill-54
-------
Rectisol process developers claim that the total sulfur content
of its CO2-rich stream may be reduced to 5 ppm, even when carbonyl
sulfide is contained in the feed gas. Yet, when Lurgi process develop-
ers designed the plant employing this system, the effluent contained
80 ppm. The total cost of the Rectisol system in the Lurgi facility
analyzed was about $55 million, or over 10 percent of the entire sys-
tem cost. Perhaps the cost of the better system, for the purer off-
gas, would have been significantly greater. It can be concluded that
the cost of achieving really low-sulfur concentrations in the CO2-rich
gas from solvent-based processes might be extremely high.
None of the licensors of the processes investigated in this study
have installed sulfur guards to eliminate the minute (ppm) quantities
of sulfur vented to the atmosphere. A combination of reasons may be
responsible for this decision:
The permissible quantities of emission of reduced sulfur
compounds have not been specified, and the effluents con-
tain less sulfur than the permissible quantities of sulfur
dioxide.
The primary sulfur compound released is carbonyl sulfide,
which is odorless; therefore, the facility will not create a
nuisance problem. Similarly, toxicity factors for car-
bonyl sulfide are not established.
The cost of maintenance of sulfur guards would be high,
because approximately 0. 9m ton (1 short ton) of iron oxide
must be regenerated or replaced during each operating day
if the off-gas contains only 50 ppm of sulfur compounds.
The applicability and design conditions for the sorption of
carbonyl sulfide on most standard sulfur guards are not
well established.
* *
The attached block flow diagrams illustrate several combinations
of acid-gas treatment and sulfur recovery. Figure III-l illustrates
the simple acid-gas treatment system proposed for the Synthane proc-
ess, followed by a Stretford process to recovery the elemental sulfur.
Figure III-2 illustrates a novel approach suggested in the Lurgi proc-
ess. In this case, the majority of the sulfur is removed in a pressure
111-55
-------
O\
©
ACID GAS
TREATMENT
REGENERATION
STRETFORD
REGENERATION
CO2,
H2S, COS
]©
GUARD
Stream Composition (Gm-Mole/Sec)
Compound
CH4
CO
C02
H2
H2O
H2S
COS
s
S02
Stream
1
1,923
1,425
3,651
4,553
21
170.0
1.95
--
•
2
1,923
1,425
11
4,553
trace
--
--'
3
3,640
-- •
.04
1.95
--
--
4
::
--
--
--
170.0
--
CO
•& §"
I—1 I—1
fl>' ^
H- C O
^3 ^^ 4n
d: £ " I-H
^S)0§
^ « m M
|1 £5
P3 2 03 t-<
P «2 H- '
» ^ B ^
W g (D
g°a
O
-------
PRESSURE
STRETFORD
ACID GAS
TREATMENT
REGENERATION
REGENERATION
C02,
H2S,
COS
Stream Composition (Gin-Mole/Sec)
on
Compound
CH4
CO
co2
H2
H2O
H2S
COS
S
SO2
Stream
1
1,923
1,425
3,651
4,553
21
170.0
1.95
--
--
2
1,923
1,425
11
4,553
--
trace
--
--
--
3
--
--
3,640
--
--
.08
1.95
--
--
4
--
--
--
' - -
--
--
--
170.0
--
-TO
2.P.
.. 4 to
T3 fl)
l-J fU
£0
tn
CO
-------
Stretford unit, and a bulk conventional acid-gas treatment (e. g., Hot
Potassium or Rectisol process) is used to recover the carbon dioxide
from the system. Figure III-3 presents the two-stage acid-gas treat-
ment system suggested in the Hygas and Texaco flow sheets. In this
case, a light acid-gas treatment is used to recover the majority of the
sulfur with minimum sulfur dioxide removal. The off-gas from this
unit goes to a Glaus plant for sulfur recovery. The second stage of
acid-gas treatment recovers the majority of the carbon dioxide.
The stream compositions for the three treatment systems illus-
trated in these figures are based upon the composition of the Hygas
system operating on high-sulfur Illinois coal. The carbonyl sulfide
concentration is approximately 1 percent of the hydrogen sulfide con-
centration. In each case, the sulfur equivalent to the hydrogen sulfide
concentration in the gas is recovered in the elemental form; however,
most of the carbonyl sulfide is discharged to the atmosphere.
Note that the COS concentration in the raw gas is based only on
thermodynamic considerations. The actual concentration in practice
may be higher; hence the quantity of COS emitted with the CO2~rich
stream may also be greater.
In Figure III-l, nearly all of the acid-gas is stripped from the
process stream. The Stretford process reduces the ^S level in the
regenerated gas to 10 ppm; the carbonyl sulfide, however, passes to
the atmosphere. In Figure III-2, the pressure Stretford unit again
reduces the hydrogen sulfide concentration of the process gas to
10 ppm, but leaves the carbonyl sulfide. These sulfur compounds
are removed in the acid-gas treatment unit and report to the CO2~
rich effluent. * Slightly better performance is achieved in Figure III-3,
where the carbonyl sulfide is divided evenly between the two acid-gas
treatment sections; however, half of the carbonyl sulfide still reports
to the CO2~rich stream and is vented to the atmosphere.
On each of these flow sheets, a "typical" acid-gas treatment
section has been assumed. Note that if a solvent-based treatment
system were used, some methane and ethane would report to the
CO2~rich stream. If a hot carbonate system were chosen, the
carbonyl sulfide would have been largely hydrolyzed to hydrogen
In a Hot Potassium acid-gas system, the COS reported in
stream 3 of Figure III-2 would have been regenerated as
but the overall sulfur emission is similar.
111-58
-------
ACID GAS
TREATMENT
I
ACID GAS
TREATMENT
II
REGENERATION
I
REGENERATION
II
HS
Stream Composition (Gm-Mole/Sec)
Compound
CH4
CO
C02
H2
H2O
H2S
COS
s
SO2
Stream
1
1,923
1,425
3,651
4,553
21
170.0
1.95
2
1,923
1,425
11
4,553
--
trace .
-'-
3
--
--
3,640
T-
--
.04
.97
-. --
--
4
170.8
5
0.2
CQ
U S
5 O o
CO
-------
sulfide. In that case, Figure III-l might have shown significantly
lower sulfur concentration in this CO2-rich gas, and Figure III-3
would have exhibited 50 ppm HgS in this effluent according to the de-
velopers of the Benfield process. The clean-fuels processes which
are in the most advanced stages of development however, do not em-
ploy hot potassium carbonate scrubbing. Perhaps undesirable side
reactions occurring in this system have discouraged its application.
If amine processes were used, the COS would form stable complexes
with amines, requiring additional reclamation. Also, amine systems
are known for high steam requirements, possibly the reason for their
not being chosen in these processes. The processes in the most ad-
vanced stages of development all use solvent-based processes for
selective acid-gas removal similar to that shown in Figure III-3. It
has therefore been assumed that the complicated flow system, incor-
porating the Rectisol and Glaus processes, is either less expensive,
more reliable, or better proved than these other alternative schemes.
These brief analyses indicate that the COS in the raw gas causes
the major contribution to sulfur emissions even though the initial COS
concentration was very low. Exxon Research and Engineering Com-
pany, in its report to the EPA on the Kopper-Totzek process (EPA-
650/2-74-009a, January 74), suggests that the COS can be catalytically
hydrolyzed to H^S. However, the initial reference quoted by Exxon
(Hydrocarbon Processing, Vol. 52, No. 2, February 73, p. 81) used
Claus-type feed gases without H2 and CO present. In the clean fuel
processes, a thermodynamic I^-CO-I^S-COS equilibrium must be
overcome to hydrogenate the COS.
4. COST ANALYSIS PROCEDURES
The costs of the pollution control and treatment schemes de-
scribed in each of the clean-fuel processes analyzed in this report
were determined by applying each of two financial accounting sys-
tems: the discounted cash flow (DCF) and utility financing methods.
Reasons for the selection of these methods and the derivation of in-
cremental.gas cost formulae for each are discussed in this section.'
It is emphasized that the costs presented in the process synopses of
Section 2 are only preliminary estimates; the actual costs of emission
control will not be known until these processing facilities have been
constructed and operated.
111-60
-------
(1) General Assumptions and Conventions
The assumptions and conventions used in the cost analysis
were as follows:
The format used for each of the financial accounting
systems is based on that proposed in the Synthetic
Gas-Coal Task Force Report to the Federal Power
Commission, dated April 1973. Specifically, the
formats used in computing capital and operating
costs for clean-fuel plants are shown in Fig-
ures III-4 and III-5, respectively.
Plant facilities were sized to generate approximately
250 x 109 Btu/day of clean-fuel energy, the daily
production expected from similar commercial plants.
Plants were assumed to operate at 90 percent of
capacity continuously throughout the year. Scheduled
plant turnarounds and emergency downtime account
for the remaining 10 percent.
The Federal income tax on gross annual income is
assumed to remain at 48 percent.
No escalation rate is assumed for operating costs
in later years; current levels of wage rates are
assumed throughout this study. If wages in the clean-
fuel industry rise more rapidly than the inflation
in the GNP, then the relative costs of pollution con-
trol computed in this report will also increase.
(2) The DCF Method
There are a number of reasons that the DCF method is an
attractive approach for computing the incremental costs of pol-
lution control in clean-fuels processes.
The DCF method of financial accounting considers
both capital and operating costs simultaneously,
using a present value approach. This minimizes
the effect of misallocating costs between investment
and operating capital requirements.
111-61
-------
FIGURE III-4
Format for Calculating Incremental *
Capital Investment
Incremental plant investment
On-site plant sections' *'
Utilities and off-site sections"'
Contractor's overhead and profit
Engineering and design costs
Subtotal plant investment
Project contingencies •
Process development contingency '*• '
Incremental plant investment (I)
Interest during construction (IDC)
Startup costs^4) (S).
Incremental working capital (W)
Incremental capital requirement (C)
Codt ($ Million).
XXX
XXX
XXX
XXX
XXX
XXX
XXX
XXX
XXX
XXX
XXX
XXX
Notes:
(1) Includes pollution abatement and recovery processes, and auxiliary
on-site investments.
(2) Including water treatment, cooling towers, power generation and distribution,
steam generation, site preparation, control houses, etc. required for pollution
controls.
(3) Not required for processes already developed
(4) Based on capitalization of 40 percent of the full-rate gross operating costs
during a six-month startup period. Assumes that 60 percent of the costs
during the startup period are covered by revenue from gas deliveries.
Incremental costs in this report refer to that portion of the total
plant cost that is assignable to pollution control facilities.
111-62
-------
FIGURE III-5
Format for Calculating Incremental
Annual Operating Cost
Labor
Cost ($ Million)
Direct operating labor
Maintenance labor
Supervision
Incremental labor cost
Administration and general overhead
XXX
XXX
XXX
XXX
XXX
Other direct costs
Raw materials
Catalysts and chemicals
Purchased utilities (electrical power,
raw water, ^' cooling water)
XXX
XXX
XXX
Supplies
Operating
Maintenance
Incremental cost of supplies
XXX
XXX
XXX
Local taxs and insurance
Incremental gross operating cost
XXX
XXX
By-product credits
Sulfur
Ammonia
Light oil
Tar
Char
by-product
XXX
XXX
XXX
XXX
XXX
XXX
Less by-product credits
Incremental net operating cost (N)
-XXX
XXX
Notes:
(1) Includes allowance for cleanup of power plant stack gases.
(2) Includes cost of delivery to plant gate.
HI-63
-------
This method permits comparisons of the incremental
costs of the various processes analyzed by discounting
all costs to the same point in time, the present.
The DCF method is widely used in industry today to
derive private investment financing requirements.
Many clean-fuel processes are expected to be devel-
oped, contracted and managed in the private sector,
and as a result the DCF method indicates the costs
these firms will incur for pollution control.
In order to apply the DCF method in the computation of incre-
mental pollution control costs for clean-fuel plants, several
assumptions had to be made.
Control equipment is funded with 100 percent equity
capital valued at 12 percent per year.
Pollution control plant facilities are depreciated
using a 16-years sum-of-digits method, based on
incremental plant investment.
Project life is 25 years.
The derivation of the formulae for calculating the incre-
mental cost of clean-fuel production due to pollution controls for
the DCF method is presented in Figure III-6. The heading of
each column explain the derivation of the results in that column.
Letting the total discounted cash flow equal zero yields the re-
quired incremental annual cost of gas, X, for the assumed
12 percent rate of return. From the figure,
4. 0784 (X-N) - 0. 971321 - 0. 52S - 0. 9412W = 0.
and solving for X yields
= 4. 0784N + 0. 971321 + 0. 52S + 0. 9412W
4.0784
or:
X = N + 0.238161 + 0. 1275S+ 0. 230777W
This equation is used to compute the incremental costs of pollu-
tion control when DCF financing is applicable.
111-64
-------
(3) This Utility Financing Method ["".'[ ' •
*•-' • ••• • • .'' ' •
This accounting method is based on a procedure developed
in 1961 by the American Gas Association. It is currently being
widely used as the primary basis for investment decisions made
by public utilities. Most of the clean-fuel processes analyzed
in this study will probably be funded in the public sector and
operated by regulated utilities using this method; .
The following assumptions were made to derive the cost
relationships for the utility financing method:. : '
Project life of 20 years. * This is reasonable for the
type of commercial plants which are likely to be built.
The incremental capital investment (less working
capital) is amortized on a straight -line basis.
', ';-; l> ''•-•' '"'•-•
Debt/equity ratio is 75 percent/25 percent.
The interest paid on debt is .assumed, to be 9 percent.
Equity funding is valued at 15
Based on these assumptions an expression cari^Se derived for
the incremental cost of product fuel resulting from pollution
control requirements. To facilitate the discussion, the follow-
ing additional terms and symbols will be used:
1 - J
n = 20 year project life - expected life-pf plant facilities
d = percent of debt capital • .
i = percent interest on debt
r = percent return on equity
Base Rate = incremental capital requirement less accrued
depreciation (includes 1/2 year, depreciation in
the given year).
The difference in project life assumed for DCF and Utility Financ-
ing reflects established accounting practice. The actual life of the
plant may differ from both of these numbers.
111-65
-------
.The average annual cost of clean-fuel is computed from the sum
of four terms which are derived as shown in Figure III-7.
The 20 year cost of clean-fuel
= the required return on capital investment = [id + r(l-d)] [20C - 10(C-W)]
+ outlay for FIT + ^ r( 1 -d) [ 20C -. 10(C-W)]
D ^
+ depreciation + (C-W)
+ net operating cost + 20N
= 20N + (C-W) + fid + r(l-d) + ^ r(l-d)l [20C- 10(C-W)]
= 20N + (C-W) + 10 [id + r( 1 -d) + ^~ r( 1 -d)] (C+W).
The average annual cost of pollution control (X) is obtained by
dividing the above expression by 20. The result is:
X = N +0.05(C-W)+ 0.5 fid + r(l-d)+ — r(l-d)] (C+W)
for i - 9 percent, r = 15 percent and d - 75 percent, one obtains:
X = N + 0.05(C-W) + 0.5 (0.105 + 0.0346) (C+W)
= N + 0.05 (C-W) + 0.0698 (C+W)
or:
X = N + 0.1198C + 0.0198W
This is the expression used in computing incremental pollution
control costs for all processes considered assuming the utility
financing method to be applicable.
(4) The Incremental Cost of Clean-Fuel
Having found the incremental annual cost of fuel (X) in
$/yr, and given the annual production rate as G (Btu/yr), the
incremental cost of fuel, due to pollution controls, is expressed
as $/106 Btu by X/G.
The relationships derived and explained above are used to
determine the incremental cost of producing each of the clean
fuels selected for detailed study in Section 2 of this report.
111-66
-------
FIGURE III-6
Derivation of DCF Equation for
Incremental Price of Clean Fuel
End Year
0
1
2
3
4
5
6
7
5
9
10
11
12
13
14
15
16
17
18
19
20
21
22
23
24
25
Totals
©
Incremental
Cost of
Clean Fuel
O
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
25X
©
Incremental N et
Operating Cost
S
N
X
N
N
N
N
N
N
N
N
N
N
N
N
N
N
N
N
N
N
N
N
N
N
N
25N+S
Depreciation
0
0.11761*
0.11031
0.10291
0.09561
0. OS82I
0. OS09I
0.07351
0.06621
0.05331
0.05151
0.04411
0.03631
0. 02941
0. 02211
0.01471
0.00741
0
0
0
0
0
0
0
0
0
I
©
©-©-©
Taxable Income
-S
X-N-0. 11761
X-N-0. 11031
X-N-0. 10291
X-N-0. 09561
X-N-0. 03321
X-N-0. OS09I
X-N-0. 07351
X-N-0. 06621
X-N-0. 05881
X-N-0. 05151
X-N-0. 04411
X-N-0. 03681
X-N-0. 02941
X-N-0. 02211
X-N-0. 01471
X-N-0. 00741
X-N
X-N
X-N
X-N
X-N
X-N
X-N
X-N
X-N
25X-25N-1-S
.52 x@
Net Income After Federal
Income Tax (FIT) of 43?o
-0. 52S
0. 52(X-N)-0. 061151
0. 52(X-N)-0. 057361
0. 52(X-N)-0. 053511
0. 52(X-N)-0. 049711
0. 52(X-N)-0. 045861
0. 52(X-N)-0. 042071
0. 52(X-N)-0. 038221
0. 52(X-N)-0. 034421
0. 52(X-N)-0. 030581
0. 52(X-N)-0. 026781
0. 52(X-N)-0. 022931
0. 52(X-N)-0. 019141
0. 52(X-N)-0. 015291
0. 52(X-N)-0. 011491
0. 52(X-N)-0. 007S4I
0. 52(X-N)-0. 003851
0. 52(X-N)
0. 52(X-N)
0. 52(X-N)
0. 52(X-N)
0. 52(X-N)
0. 52(X-N)
0. 52(X-N)
0. 52(X-N)
0. 52(X-N)
13(X-N)-0. 521-0. 52S
1 ©
Investment*
1. 23676I-W
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
-w
1.236761
©' © - ©
Cash Flow
-1. 236761-0. 52S-W
0. 52(X-N)+0. 0564SI
0. 52(X-N)*0. 052941
0. 52(X-N>+0. 049391
0. 52(X-N) +0. 04589!
0. 52(X-N)tO. 042341
0. 52(X-N)-0. 033831
0. 52(X-N)*-0. 035281
0. 52(X-N)+0. 0317SI
0. 52(X-N)*0. 028221
0. 52(X-N>0. 024721
0. 52(X-N)^0. 021171
0. 52(X-N)tO. 017661
0. 52(X-N)-0. 014111
0. 52(X-N)»0. 010611
0. 52(X-N)tO. 007061
0. 52(X-N)*0. 003551
0. 52(X-N)
0. 52(X-N)
0. 52{X-N)
0. 52{X-N)
0. 52(X-N)
0. 52(X-N)
0. 52(X-N)
0. 52(X-N)
0. 52(X-N) +W
13(X-N)-0. 75681-0. 52S
©
& 127°
Discount Factor
1.0000
0.8929
0. 7972
0.7118
0. 6355
0. 5674
0. 5066
0. 4523
0. 4039
0. 3606
0. 3220
0.2875
0.2567
0.2292
0. 2046
0. 1827
0.1631
0. 1456
0. 1300
0.1161
0. 1031
0.0926
0.0826
0. 0738
0. 0659
0.0588
© * ®
Discounted Cash Flow
-1. 236761-0. 52S-W
0. 46431(X-NV-0. 050401
0. 41454(X-N)+0. 042201
0. 37014(X-N)+0. 035161
0. 33046(X-N>tO. 029161
0. 29505(X-N)+0. 024021
0. 26343(X-N)+0. 019671
0. 23520(X-NKO. 015961
0.21003(X-N)*0. 012841
0. 18751(X-N>"-0. 010181
0. 16744{X-NHO. 007961
0. 14950(X-N)+0. 006091
0. 13348(X-N)+0. 004531
0. 11918(X-N)-!-0. 003231
0. '0639(X-N)+0. 002171
0. 09500(X-N)+0. 001291
0. 08481(X-N)+0. 000581
0. 07S71(X-N)
0. 06760(X-N)
0. 06031(X-N)
0. 05392(X-N)
0. 04815(X-N)
0. 04295(X-N)
0. 03838(X-N)
0. 03427(X-N)
0. 03058(X-N) +0. 0588W
4. 07840(X-N)-0. 971321-0. 52S-0. 9412W
-*Sum-of-years-digits depreciation of the incremental plant
n + 1 - age of asset , , . . ,
investment is n (n + l)/2 where n = depreciated 5
life. For the first year it is 16/(16(17)/2) I = 0. 11761.
^Capitalized costs in year 0 include incremental plant investment
(I), return on investment during construction, and incremental
working capital (W).
LEGEND
S Incremental startup costs
X Incremental annual cost of fuel
N Incremental net operating costs
I Incremental plant investment
W Incremental working capital
-------
FIGURE III-7
Derivation of Utility Financing Equation
For Incremental Cost of Clean Fuel
End Year
1
:
y**
Sum of v years
Sum of 20 years
(y = m = 20 years)
(A)
Required Rate of
Return on Investment*
id + r ( 1 - d 1
id + r (I - d )
1
1
1
I
;
id + r 11 -d)
©
Depreciation*
1
— (C • W)
in
1
— (C - W)
m
i
I
I
I
I
- IC-W)
m
y
- (C-W)
m
C-W
©
Accrued
Depreciation
' fC"W\
- \ m 1
3 /C-W\
~(~ 1
1 i\/c-w\
Hb)
©
Capital Investment -(C^
Rate Base
1 /C-'W \
2 \ m /
1
1.
1
1
I
1 'Vc-w\
c-b)U)
©
(A)X(D)
Return on Capital Investment
r ir -/c-Mi
.d + r (1 -dl C- • (
L JL 2V m i\
[- IF ,/c.w\l
L JL 2\ m /J
i
i
i
i
i
[—][-(> -i)(c^)]
r ll xf i\/c-«\l)
I id + r(l • d) |)yC - J |[y — V 1 1 i
L J| 'L\ 2Am ''Ji
id *t (1 -d) 20C- 10(C-W)I '
©
Cost of Equity Funding x^Dj
Return on Equity
[1 / C - w \ ]
2V m )\
[3 /C-W \]
C-- ( I
2 \ m /J
r(l-d) c-(y--)(— I
[ \ 2/\ m /J
©
48. x(9)
52 ^-^
Federal Income Tax (F.I.T.)
48 r i /c-w\
52 [ 2 \ m /j
48 f 3 /C-W \
— id -d) C-- 1 I
52 |_ 2 \ ra /J
1
1
1
1
1
48 f / 1\/C-W\"|
- rd -d) C- y--)[
52 [ V lA m /)
48
— i(I - d)
52
r lN/CtW\"|l
yc- 2 \ly--\~- 1
1 |\ 2^1" /J I
r i
— r d -d) 20C-10 (C-W)I
52 L J
®
Net Operating Cost*
N
N
1
1
1
1
N
yN
20N
00
Weighted average of cost of debt and equity funding respectively
I Straight-line depreciation including —years depreciation for given year
^ Excludes escalation of operating costs in future years
*" Any given year
-------
SECTION 2
IV. THE SYNTHANE PROCESS
The Synthane process has been under development by the U.S.
Bureau of Mines (USBM) since 1963, at the Pittsburgh Energy Research
Center in Bruceton (14 miles south of Pittsburgh). The process is
designed for the conversion of bituminous and sub-bituminous coal
and lignite into sulfur-free substitute natural gas (SNG). The primary
effort reported here has been directed at the 1. 6 percent sulfur
Pittsburgh seam coal and the 2. 9 percent sulfur Illinois No. 6 coal.
The process can be modified to use air instead of oxygen to produce
a low Btu gas.
The first USBM unit was evaluated by the M. W. Kellogg Com-
pany; Hydrocarbon Research, Inc., was engaged to develop a com-
mercial process design. Designs for a 68 m ton/day (75 short tons/
day) pilot plant to be constructed in Bruceton were completed by the
Lummus Company in 1972, and the construction contract has been
awarded to the Rust Engineering Company.
The purpose of the $14 million plant will be to provide informa-
tion on technical feasibility and costs of a commercial installation.
It will feature a 71. 3 kg/cm2 (1000 psi), 1. 2 m ID gasifier capable of
generating 34, 000 m3/day (1.2 x 10° ft /day) of gas. Completion and
start-up tests are scheduled for September 1974.
1. PROCESS DISPLAY
The Synthane process has recently been modified to include ex-
ternal pretreatment of the coal. This modification was too new for
information to be available for this process synopsis.
The flow diagram and tabulations of associated mass and
energy balances for the prototype Synthane coal gasification process
are presented in Figure IV-1.
IV-1
-------
(1) Basis for Analysis
Process display is for a 1.6 percent sulfur
Pittsburgh seam ccJal, but estimated sulfur
emissions from a 2. 9 percent sulfur Illinois coal
are also reported, where available.
Production rate is 7.08 x 106 m3/day (250 x 106 ft3/
day) of gas, containing 58. 41 x 109 kcal (231. 8 x
10^ Btu), from 12, 930 m tons/day (14, 250 tons/day)
of coal, containing 91.22 x 109 kcal (362.0 x 109
Btu), yielding a coal-to-gas thermal efficiency of
0.640.
(2) Layout and Symbols
The general direction of process flow is from the coal
and water supply on the left to the product gas on the extreme
right of the flow diagram, with residuals shown along the
bottom. The bold line'indicates the flow of the primary gasifi-
cation process. Portions of the flow diagram enclosed within
the dashed boundary were specified by the U.S. Bureau of Mines;
the remaining steps are proposed for the present analysis.
The figures in circles and triangles refer to .the stream
composition tables in the lower portion of the figure. Figures
in brackets report the sulfur flow rate in short tons per hour,
for the 1.6 and 2.9 percent sulfur, respectively. Other figures
and tables are self-explanatory. The overall energy balance is
shown to the right of the material balance tables.
Rhombic-shaped units represent intermediate products,
uses, and sources, for which the distribution is not shown.
Nonintegral pollutant cleanup processes are indicated by
sloping rectangles. The treatment and recovery processes,
represented here are described in the process description and
pollution control sections of this process summary. Inverted
trapezoids denote residual storage.
2. PROCESS DESCRIPTION
The specific steps of the Synthane coal gasification process are
described in this section, and the generation of pollutants is discussed
-------
FIGURE IV- 1
The Synthane Process
CO'J'O^ENT
^
2
' H3
CM4
FLOW RATE
±-»x;^
COMPOSITION OF CASEOUS STREAMS 'VOL. V
G>©©0©®0®®® ®®®©@®
'
1
MB; 0.76S 0X2 O.IIO «09] O834 0791 O.U8 OBIT 1.3V9 1.I1B DM2 O.I9S 0.263 1135 MM
01 ISO 3»
COMPOS OF LIQUIDSTREAMS (me/I!
COMPONENT
SUSP. SOt IDS
COD
CO,
NH,
PHENOL
CYANIDE
THIOCYANATE
S
FLOW Itm
RATE gpm
A A &
33
19.000
1WCO 13JXD 13.000
11,000
UOD
OJt
ISO -
300 300 300
i!na ?>• i*s7
A
a
•.707
13 JOB
3J»
603
0.2
09
300
14J49
3.791
CONSTIT.
c
H
o
N
s
H,0
ASH
FLOW
RATE vfc
SOLID AJ
^
n»
61
«.o
1.5
1.6
2»
7.4
893.71
WALVSIXW
^
71.4
0»
I*
OS
l.S
-
23.9
181.10
rr.%j
^
774
7.0
7.3
-
S3
-
aj
M.9Q
CARRIER
COAL
TOT At. INPUT .
PRODUCT GAS
TAfl
SULFUR
AMMONIA
COOLING WATER
OTHER at DIFFERENCE)
. TOTAL OUTPUT
ENERGY BALANCE
ENERGY (10* kc«mr>
J-tOCLS
.. JJQft»
I.433.S
100.7
11 J
17.9
OIU
1»«.4
ijoo.»
ENERGY HO* BUM
1S.OSU
tljOHJ.
9«aa.7
•n.7
44.4
71.0
1X10.0
771 .»
tjjjji^
-------
where appropriate. Pollutant control and cleanup processes are
covered in subsequent sections.
(1) Coal Preparation
In the coal preparation stage, the coal is crushed by
hammer mills and ground by ball mills until 70 percent of it
passes through a 200-mesh screen (average diameter of 95
microns). The pulverized coal is then conveyed by an air
stream to the lock hoppers. Coal dust and runoff wastewater
are generated during this operation.
(2) Gasification
Coal from the lock hoppers is fed into the top of the fluid-
bed gasifier, while a mixture of oxygen and steam is introduced
at the bottom. The first stage of the gasifier provides pre-
treatment of coal in the form of mild oxidation at 430 degrees C
(800 degrees F) to prevent agglomeration. The gases formed
here, including methane, carbon monoxide, and hydrogen,
become part of the raw gas. The process does not require
pretreatment that produces a separate waste stream.
In the lower stages, pulverized coal, steam, and oxygen
react at 980 degrees C (1800 degrees F) and 35-70 kg/cm2
(500-1000 psi) to produce raw gas, char, and tar. Char residue
is withdrawn at the bottom of the gasifier, while the raw gas is
withdrawn. The quantities of major constituents are shown in
Figure IV—1. Recent process development unit (PDU) data also
indicate the presence of suspended solids, COD, phenol, am-
monia, cyanide, thiocyanate, carbonyl sulfide, thiophene, and
methyl mercaptan. These compounds are removed during
subsequent purification.
(3) Raw Gas Cleanup
The raw gas from the gasifier is first passed through
cyclone separators to remove entrained char particles and then
is scrubbed with water and oil to remove tar and other impu-
rities. The composition of the effluent gas stream is shown in
Figure IV-1» while the breakdown of sulfur compounds present
is reported in Table IV-1.
IV-4
-------
Table IV-1
Sulfur Compounds in Scrubber Effluent Gas
Component
Carbonyl sulfide
Thiophene
Methyl mercaptan
Hydrogen sulfide
ppm
! (by volume)
11
55
8
(by difference)
Tons/Day
(Sulfur)
0.5
2.5
0.3
131.1
The scrubber water is recycled through a decanter that
separates the tar and bleeds off excess wastewater, with the
composition given in Table IV-2 and shown as stream /r\ in
Figure IV-1. The composition of the tar appears in Figure IV-1
under column^1, but recent PDU data for Illinois No. 6 coal
indicate also the presence of trace quantitites of arsenic
(0. 7 ppm) and mercury (0. 003 ppm).
Table IV-2
Composition of Wastewater From Gasifier
Component
Suspended solids
COD (chemical oxygen demand)
Carbon dioxide
Ammonia
Phenol
Cyanide
Thiocyanate
Sulfur
Aromatics
A.
Pyrfdines
Mg/Liter
23
19,000
13,000
11,000
1,700
0.6
188
300 ppm
trace
trace
Lbs/Hr
15
'12,710
8,680
7,360
1,140
<1
130
200
~
~
IV-5
-------
(4) Shift Conversion
~!
following the scrubbing state, the gas stream undergoes
CO shift conversion to adjust the H2/CO ratio to approximately
3.2 prior to the purification and methanation stages. The gas
is first] split into two streams. The first stream is heated to
430 degrees C (800 degrees F) by hot gas from the shift reactor
and is fed, along with steam, to the reactor. The shifted gas
is then mixed with the second half of the scrubber stream and
passed through a waste heat recovery boiler (WHR) to recover
heat. The composition of the condensate produced here is
shown as stream £^ in Figure IV-1 under the assumption that
the sulfur and carbon dioxide content is the same as in the
raw gas condensate.
(5) Purification
; The acid gas purification stage selected by the USBM
consists of single-stage scrubbing by a hot potassium carbonate
solution to remove carbon dioxide, hydrogen sulfide, and car-
bonyl sjulfide followed by cooling, passage through a knock-out
drum, and then through iron oxide and char towers to remove
any remaining traces of hydrogen sulfide and other sulfur com-
pounds ^that would poison the methanation catalysts. About
30 pprn of hydrogen sulfide and carbonyl sulfide, and 40 ppm
of thiophene and mercaptans are expected to be present in the
gas stream leaving the scrubber.
The spent hot potassium carbonate solution is regenerated
by steam to produce a sulfur recovery feed stream containing
about 1.5 percent hydrogen sulfide and a condensate with a
composition shown as stream & in Figure IV-1. Loss of hot
potassium carbonate solution is minimal and is not expected
to present a pollution problem.
Spent iron oxide is discarded or returned to the manu-
facturer for regeneration, while activated carbon is regenerated
by steam. After condensation of steam, the remaining substance
can be incinerated. The amounts of emissions produced here
are quite small and are not expected to contribute significantly
to the pollution problem.
IV-6
-------
(6) Methanation
The purified gas is methanated over Raney nickel catalyst.
Its amount and life are not known, but the catalyst will be re-
turned to the manufacturer for reactivation or replacement. In
the commercial operation, hot gas recycle (HGR) methanation
will be substituted for the original concept using Dowtherm, a
high-temperature liquid which could decompose and contribute
to the pollution problem. The HGR method utilizes the sensible
heat capacity of the recycled gas to cool the catalyst.
(7) Cooling and Drying
Following methanation, the product gas is cooled to 38
degrees C (100 degrees F) to condense out water, which contains
very little contamination and can be reused. Although in the
original flow diagram there was no provision for drying the gas
to pipeline specification of 7 pounds of water per 10 ft', of gas
(112 kg of water per 10° m^ of gas), this can be accomplished
with the aid of a drying system using ethylene glycol. The
amounts of water and solvent given off at this stage are quite
small and are not expected to add significantly to the pollution
problem.
(8) Steam and Power Generation
The steam and electric power requirements of the process
are to be met by an on-site steam plant fired by the gasifier char.
The resultant stack gases would contain sulfur dioxide in excess
of the maximum permissible level of 1.2 pounds of SC>2 per 10°
Btu of fuel burned (2.16 kg/106 kcal). The heating value of the
char supplied for steam and power generation totals 994 x 106kcal/hr
(3940 x 10° Btu/hr). The process developer reports a power
requirement of 6000 kW. Based on the projections of other de-
velopers, 40 to 50 MW seems more realistic. The estimated
total high pressure steam requirements for this facility are
2, 478, 600 Ib/hr (1, 124, 590 kg/hr). Sufficient energy is avail-
able from the char for this steam and power requirement. '
The oxygen requirements could be provided by an on-site
steam-powered oxygen plant, which produces no harmful emis-
sions.
IV-7
-------
(9) Energy Balance
The overall energy balance for the Synthane process is
calculated from the heats of combustion and sensible heats of
the feedstocks, products, and residuals. A summary tabulation
of such calculations is presented in Figure IV-1 and derived
in Table IV-3 on the basis of heats of combustion listed at the
beginning of this chapter. The energy conversion efficiency
(coal-to-gas) is, therefore:
2.433.5/3,800.5 - 0.640.
(10) Sulfur Balance
Sulfur flow rates for the various steps in the process
have been calculated for two different feedstocks (1.6 percent
S Pittsburgh seam coal and 2. 9 percent S Illinois No. 6 coal)
and for two alternative methods of char treatment (direct com-
bustion and gasification plus combustion). Sulfur content of
various streams was obtained by extrapolation from process-
development unit data or from related commercial experience.
The resultant flow rates, expressed as m tons/hr and short
tons/hr of sulfur are indicated in Figure IV-1 and tabulated in
Table IV-4 of this process summary.
The reported sulfur output is the total evolved from all
sources. It is composed of sulfur remaining in the by-product
tar which is sold, sulfur recovered in its elemental form. Sul-
fur is also released to the atmosphere as SOn from the iron
oxide/char towers. Not all of the sulfur in the feed stream can
be recovered by the Wellman-Lord and Stretford processes.
(11) Water Requirements
"The process developer projects a water requirement of
56,000 liters per minute (15, 000 gpm). This indicates exten-
sive use of water cooling within the process. Based upon the
expectations of other developers and the FPC Task Force Re-
port (Ref. 4) this water load might be reduced by a factor of
two or three if extensive use is made of air cooling.
IV-8
-------
Table IV-3
Energy Balance Calculations
Carrier
Coal
Total energy input
Product Gas
Tar
Sulfur
Ammonia
Cooling Water
Other*
Total energy output
Calculation
538,650 kg/hr x 7,056 kcal/kg
295,009 m3/hr x 8,249 kcal/m3
18,960 kg/hr x 8,476 kcal/kg
5,069 kg/hr x 2,2 13 kcal/kg
3,338 kg/hr x 5,372 kcal/kg
984,100 1/min x 59.9 kg min/1-hrs x 16.7 kcal/kg
(by difference)
Energy (106kcal/hr)
3,800.5
3,800.5
2,433.5
160.7
11.2
17.9
981.8
194.4
3,800.5
(106Btu/hr)
15,081.3
15,081.3
9,656.7
637.7
44.4
71.0
3,900.0
771.5
15,081.3
* Includes sensible heat of product streams, heating value of other products,
and heat lost to the atmosphere.
Table IV-4
Sulfur Balance - Short Tons/Hr (m Tons/Hr)
Carrier
Coal
Total input
Tar
Sulfur
Sulfur (emissions)
Sulfides from ,
sulfur recovery
Sulfides from
iron oxide and
char towers
Total output
Pittsburgh Seam Coal
Direct
Combustion
9.50(8.62)
9.50(8.62)
1.14(1.03)
7.07 (6.42)
1.18(1.07)
0.06 (0.05)
0.05 (0.05)
9.50 (8.62)
Indirect
Combustion
9.50(8.62)
9.50(8.62)
1.14(1.03)
8.13(7.38)
0.10(0.09)
0.08 (0.07)
0.05 (0.05)
9.50 (8.62)
Illinois No. 6 Coal
Direct
Combustion
20.17(18.30)
20.17(18.30)
0.28 ( 0.25)
18.41 (16.79)
1.25 ( 1.13)
0.18 ( 0.16)
0.05 ( 0.05)
20.17(18.30)
Indirect
Combustion
20.17(18.30)
20.17(18.30)
0.28 ( 0.25)
19.54(17.73)
0.10 ( 0.09)
0.20 ( 0.18)
0.05 ( 0.05)
20.17(18.30)
IV-9
-------
3.
DISCUSSION OF CONTROL PROCESSES
The five major waste streams emitted by the Synthane process
and the proposed control and treatment methods are summarized in
Table IV-5.
Table IV-5
Nature and Treatment of Major Waste Streams
Final Wastes
Coal dust
to atmosphere
Char
Tar
Wastewater
Sulfur dioxide to
atmosphere
Sources
Coal pulverization and
handling
Gasifier
Scrubber
Coal storage
Scrubber
Shift converter
Regenerator
Power plant, Stretford,
iron oxide and char towers
Treatment
Cyclone separators,
bag filters, enclosure
Various
Various
Biological, Pheno-
solvan and modified
Chevron
Wellman-Lord
(1) Control of Coal Dust
Most of the coal dust produced during the coal pulveri-
zation and handling operations can be controlled with the aid
of cyclone separators and bag filters. Assuming that 5 percent
of the gross coal feed, or 27, 000 kg/hr (60, 000 Ibs/hr), gets
by the cyclone separators and that the bag filters retain 99. 9
percent of that amount, some 27 kg/hr (60 Ibs/hr) of fine coal
dust is expected to escape to the atmosphere, unless additional
precautions are taken, such as enclosure of the coal preparation
operation.
(2) Utilization of Char
The char residue withdrawn from the bottom of the
gasifier still contains approximately 70 percent of carbon,
which can be converted to energy in the on-site steam plant.
The stack gas produced, if the char were burned directly,
IV-10
-------
contains sulfur dioxide in excess of the maximum permissible
level of 2. 16 kg SO2 per 106 kcal of fuel (1. 2 lb/106 Btu), so that
some form of desulfurization is required. This can be achieved
through:
Char pretreatment
Combustion modification
Stack gas cleanup.
The type of pretreatment is determined primarily by the
nature of sulfur compounds present (for example, organic
versus pyritic). In the Synthane process, since the iron pyrite
is probably decomposed, two potentially applicable techniques
for desulfurizing raw coal or char (IGT flash desulfurization
and TRW pyrite leaching) would probably not be satisfactory.
The char could, however, be converted to a low Btu gas of about
2670 kcal/m3 (300 Btu/ft ). This gas could be then desulfurized
by acid gas cleanup similar to the scheme employed in the main
process and burned in the steam plant. This operation should
be carried out in a reactor capable of producing a clean ash—
for example, through high temperature gasification of all carbon
and sulfur. Suitable units would include:
Koppers-Totzek reactor
Bottom stage of Bi-Gas reactor
Hygas steam-oxygen reactor
U-Gas steam-air reactor
Lurgi steam-air reactor.
The sulfur balance for the overall process incorporating the
char gasification techniques is reported in Table IV-4 under
the heading of indirect combustion.
Sulfur removal during combustion could be accomplished
by the lime bed technique similar to the present EXXON-EPA
study, but this approach is not yet proved. Stack gas cleanup
was assumed for this analysis, and the sulfur balance for the
overall process incorporating this last char treatment technique
is reported in Figure IV-1 and in Table IV-4, under direct
combustion.
IV-11
-------
(3) Utilization of Tar
The Synthane process generates a relatively large amount
of tar because of the countercurrent flow of raw coal and gases
in the upper stages of the gasifier. This tar might be utilized
as an ingredient in road paving or roofing materials, or as fuel.
The latter application requires removal of the high sulfur con-
tent, as was the case with the char, and can be implemented
in one of four ways:
Conversion to oil
Direct gasification
Direct combustion
. Indirect combustion.
Direct conversion to oil would require the application of
residuum hydrocracking and desulfurization processes, such
as the Gulf HDS process. Tar from gasification processes may
not be suitable for standard distillate hydrodesulfurization
because of its probable high particulate content. Alternately,
the tar could be fed back into a low temperature zone of the
gasifier where it can be hydrocracked. The resulting oil might
be suitable for distillate hydrodesulfurization techniques.
Direct gasification could be achieved by reintroducing the
tar into the high-temperature zone of the gasifier where the acid
and gas cleanup unit would remove the sulfur from the gas. The
Bureau of Mines is currently investigating this approach.
In the case of combustion to supplement process energy
requirements, sulfur might be removed in the furnace, on a
bed of lime, or by stack gas cleanup. The latter approach is
considered demonstrated by the EPA, whereas the former is
not.
Indirect combustion might be accomplished by partial
oxidation of the tar (e. g., by the Texaco Partial Oxidation
process), followed by desulfurization and combustion of the
resultant low-Btu fuel gas. Alternatively, the tar could be
impregnated on the char and gasified jointly by one of the
techniques described in the preceding section.
IV-12
-------
(4) Liquid Waste Treatment
The principal liquid -»vaste streams are the leachings from
coal storage and the three condensate streams from the gasifier/
scrubber, the shift converter, and the regenerator. The com-
positions of the three individual condensate streams are reported
in Figure IV-1 in columns &, £^,£$± respectively, whereas the
composition of the merged condensate stream is shown in
column^ The composition of stream/^ (H2S in wastewater)
was confirmed by the process developer. Since H^S concen-
trations in water are dependent on its solubility in water,
levels were assumed to remain the same for streams
Phenol may be recovered and sold as a by-product, de-
stroyed by biological oxidation, or recycled to the reactor. In
the latter case, it must still be removed from the wastewater
stream to prevent excessive cooling and contamination of the
reactor by the other components. The Phenosolvan process
has been demonstrated commercially to remove phenol from
wastewater leaving a residual concentration as low as 20 ppm.
Some processes used by oil refineries may apply here, and
additional processes are under development. The Phenosolvan
process also removes most other hydrocarbons that may be
present.
Ammonia, carbon dioxide, and hydrogen sulfide are
recovered by the Chevron wastewater treatment process.
Hydrogen sulfide and some carbon dioxide are first removed
by a stripping tower and then sent to the sulfur recovery stage.
Ammonia is recovered by a second stripping tower for sale as
a by-product of a purity adequate for manufacture of anhydrous
fertilizer. The Phosam process developed by U.S. Steel for
recovery of ammonia from coking operations may be applicable
here as well.
The wastewater purified by the Phenosolvan and Chevron/
processes may still contain 20 to 50 mg/1 of phenol, several
hundred mg/1 of oils, and small amounts of ammonia, cyanide,
and sulfur compounds. Its quality is satisfactory for cooling
water makeup, though not for boiler water makeup. The low-
concentration contaminants in cooling water are eventually
eliminated by biological action in the cooling system, or they
wind up in the cooling tower blowdown. The latter can be used
IV-13
-------
in such applications as quenching of hot ashes discharged from
the bottom of the gasifier or the steam plant furnace.
Blowdown from cooling towers and boilers, as well as
sanitary sewage and runoff water, can be treated by a miscel-
laneous treatment section incorporating a sewage treatment
plant. The only waste product discharged from the wastewater
treatment stage will be an inert solid residue. The quantity of
this solid residue cannot be estimated since it is dependent on
the particular unit processes selected to treat this wastewater
and the method used to handle the resulting solid residue. Some
water-soluble materials (such as chromium compounds used for
boiler feed water treatment) may not be removed by this treat-
ment, thus requiring a small bleedstream to evaporation ponds.
Two possible alternatives to the above recovery processes
are high temperature, about 1400 degrees C (2500 degrees F),
incineration and biological treatment of the liquid waste. The
first method is costly because of the high fuel requirements,
whereas the second would destroy possible valuable by-products.
(5) Sulfur Recovery
Recovery of elemental sulfur from the acid gas (CO2 and
H^S) stream leaving the absorption stage is accomplished here
by the Stretford process. This process was selected in prefer-
ence to the more popular Glaus unit, because it operates more
efficiently with the low (1.5 percent) hydrogen sulfide concen-
trations prevailing here. The design efficiency of a Stretford
process is 99 percent. Even higher efficiencies may be attained
when operating on the effluent from a hot carbonate process
since f^S (easily handled by a Stretford unit) is the only sulfur
compound present. The carbon dioxide effluent from a Stretford
unit, using conventional design specifications, contains very
small amounts of sulfur components (<150 ppm). Even lower
sulfur content may be possible in this instance.
As an alternative to the Stretford process, the off-gas
from the single-state hot potassium carbonate acid gas treat-
ment could be treated by a secondary acid gas cleanup unit that
is selective for H2S. Suitable units would include amine systems,
such as tertiary amine (TEA), diisopropanol amine (DIPA), or
IV-14
-------
solvent based systems, such as M-Pyrol, Rectisol, and Selexol.
The sulfur-rich gas fraction could then be fed to a Glaus unit,
while the sulfur-lean remainder can be incinerated but at signi-
ficant energy cost due to the additional coal requirements to
fire the incinerators. Several techniques are available to treat
the Glaus tail gas.
A third alternative would be provided by the replacement
of the single-stage hot carbonate treatment by a two-stage
arrangement. Here the off-gas from the first stage would be
treated by a two or three-stage Glaus unit and its tail gas could
be treated by available techniques. The off-gas from the second-
stage hot carbonate treatment might contain 50-150 ppm HoS
which could be trapped in an iron-sponge bed again at relatively
high cost.
The above discussion is intended merely to give the flavor
of the great variety of schemes available for sulfur recovery.
A complete definition of technical feasibility and costs and a
thorough tradeoff analysis would be required prior to making a
final selection.
(6) Stack Gas Cleanup
Sulfur dioxide in the stack gases from the steam and
power plant can be recovered by lime scrubbing processes
which have been demonstrated in several related applications.
For this process analysis, we included the less proved but more
advanced Wellman-Lord system. It is expected that this system
will be proved by the time the Synthane process is ready for
demonstration. The Wellman-Lord process regenerates
concentrated SCv To arrive at the sulfur recovered in
Figure IV-1, the incorporation of an Allied process was
assumed.
A potential alternative is presented by the citrate
process being developed by the Bureau of Mines, which
is based on a reaction between sulfur dioxide and hydro-
gen sulfide. Other techniques in various stages of de-
velopment include catalytic oxidation, and the double
alkali processes.
IV-15
-------
4. COSTS OF POLLUTION CONTROL *
The costs of the pollution control and treatment processes
described in the previous section are calculated in this section. It
should be noted that the Synthane process is not currently commercial,
so that figures given in this section are based largely on extrapolations
from PDU data and related commercial experience.
The major assumptions and conventions adopted here are dis-
cussed in Chapter III, Section 1. All by-product flow rate data pre-
sented in Figure IV-1 apply.
The incremental** capital costs for pollution control are
reported in Table IV-6 for each feedstock considered, and for
both the direct and indirect combustion of residual char for
steam and power generation. The incremental annual operating
costs for pollution control are presented in Table IV-7.
The derivation of the formulas for calculating the incremental
cost of gas production due to pollution control is presented in
Chapter III, Section 1. For the discounted cash flow (DCF) method,
the required incremental annual cost of gas, X, for the assumed rate
of return is:
X = 0.23816 I -1- 0. 1275S + 0.230777W
where:
N = incremental net operating cost
I = incremental plant investment
S = startup costs
W = incremental working capital.
For the utility financing case (UFC), it is:
X = N + 0. 1198 C + 0. 0198 W
* The costs presented here are preliminary estimates; the actual
costs of emission control in this process will not be known until
a full-scale facility has been constructed and operated.
** Incremental costs in this report refer to that portion of the total
plant cost that is assignable to pollution control facilities.
IV-16
-------
Table IV-6
_ o^
Incremental Capital Investment ($ Million)
Item
Incremental Plant Investment*5
Dust Emission Control
Tar Utilization*
Wastewater Treatment
Wellman-Lord SO-. Recovery
Allied Sulfur Recovery from SO-,
Stretford Sulfur Recovery
IGTU-Gas Indirect Combustion of Char
Subtotal Plant Investment
Project Contingencies0
Incremental Plant Investment (I)
Startup Costsd (S)
Interest During Constructione (IDC)
Incremental Working Capital (W)
Incremental Capital Investment (C)
1.67. Sulfur Coal
Direct Combustion of Char
-> -i
10.8
7.7
4.0
3.6
28.3
4.2
32.5 32.5
1.5 1.5
Utility DCF
5.5 7.7
0.7 0.9
40.2 42.6
Indirect Combustion of Char
i -,
10.8
3.6
24.9
41.5
6.2
47.7 47.7
1.8 1.8
Utility DCF
8.0 1 1 .3
1.0 1.2
58.5 62.0
2.9% Sulfur Coal
Direct Combustion of Char
2.2
10.8
6.6
3.5
8.2
31.3
4.7
36.0 36.0
2.0 2.0
Utility DCF
6.1 8.5
0.8 1.0
44.9 47.5
Indirect Combustion of Char
2.2
10.8
8.2
24.9
46.1
6.9
53.0 53.0
2.2 2.2
Utility DCF
8.9 12.5
1.2 1.4
65.3 69.1
Notes:
a Incremental plant investment, return on investment during construction,
and working capital are treated as capital costs in year 0 (the year ending
with completion of startup operation).
b. Installed costs, including engineering design costs, contractors' profit, and
overhead.
c Includes costs for unexpected site preparation and hardware requirements
at 15 percent of plant investment.
d At 20 percent of incremental gross operating cost.
e For the DCF method, computed as the discount rate x Incremental Plant
Investment for 1.875 years' average construction period.
1(1+0" = 1(1.12)1-875 = 1.236761 or 0.2367611 additional investment
For the Utility Financing method, computed as the interest rate on debt x
Incremental Plant Investment x 1.875 years.
f Sum of materials and supplies at .9 percent of incremental plant investment
and net receivables at 1/24 of annual incremental revenue.
* It is assumed that recovered tar can be fed to the gasifier and transformed
into product gas at no additional cost.
-------
Table IV-7
Incremental Annual Operating Cost ($)
Labor
Direct Operating Labor (3 men/shift x
S5/hr x 8,304 shift-hrs/man-yr)
Maintenance Labor (1.5% of I)
Supervision ( 1 5% of direct operating
and maintenance labor)
Incremental Labor Cost
Administration and General Overhead
(60% of Incremental Labor)
Other Direct Costs*
Supplies
Operating (30% of Direct Operating
Labor)
Maintenance (1.5% of I)
Incremental Cost of Supplies
Local Taxes and Insurance (2.7% of 1)
Incremental Gross Operating Cost
By-Product Credits
Sulfur (short tons/day x 365 days/yr
x 0.9 capacity factor x $ 10/long ton x
2000/2240 long tons/short ton)
Ammonia (88.3 short tons/day for 1.6%
sulfur coal and 1 10 short tons/day for
2.9% sulfur coal @ 0.9 capacity factor
x 365 days/yr x $25/short ton)
Less By-Product Credits
Incremental Net Operating Cost (N)
1.6% Su
Direct Combustion of Char
1 24.600
487,500
91,800
703.900
422.300
5.000.000
37,400
487.500
524.900
877.500
7.528,600
497,700
725,200
-1.222,900
6,305,700
fur Coal
Indirect Combustion of Char
1 24,600
715,500
1 26.000
966,100
579.700
5,200,000
37,400
715.500
752,900
1,287,900
8,786,600
572,200
725,200
-1,297,400
7,489,200
2.9% Sulf
Direct Combustion of Char
124,600
540,000
99,700
764,300
458.600
7,000,000
37,400
540,000
577,400
972,000
9,772,300
1,294,900
903,400
-2,198,300
7,574,000
urCoal
Indirect Combustion of Char
124,600
795,000
137,900
1,057,500
634,500
7,200,000
37,400
795,000
832.400
1,431,000
11,155,400
1,374,400
903,400
-2,277,800
8,877,600
<
OO
This item includes utilities and materials expended in the processes. The cost shown is a rough estimate that will need to be defined more precisely.
-------
where:
C = required capital investment
N = defined above
W = defined above.
The annual gas production is:
(250 x 106 ft3/day) (927. 1 Btu/ft3) (365 days/yr)
(0. 9 capacity factor) = 76.138 x 1012 Btu/yr or 19.198 x 1012
kcal/yr.
This is about the production rate one would expect from a commercial
Synthane plant.
The incremental cost of gas due to pollution controls is found as:
annual cost _ X
annual gas production ~ ?6< 1JJ8 x Ifll2 Btu/yr
Solutions to the above equation for the incremental cost of gas
are given in Table IV-8. The major impact on these costs, for the
Synthane process, is net operating costs (N) which comprise nearly
60 percent of the incremental cost of gas. A graphic representation
of the increase in the price of gas as a function of sulfur content of
feedstock and of incremental capital cost is shown in Figure IV-2
(by-product credits are excluded).
5. REFERENCES
(1) "An Economic Evaluation of Synthane Gasification, "
U.S. Bureau of Mines Report No. 72-9A, October 1971.
(2) "Synthane Coal Gasification Pilot Plant, " Final Environ-
^ mental Statement, U.S. Department of Interior FES 72-78.
(3) Forney, A.J., Private Communication, U.S. Bureau of
Mines, Pittsburgh Energy Research Center,
September 1973.
(4) "The Supply Technical Advisory Task Force - Synthetic
Gas Coal, " Final Report, Federal Power Commission,
April 1973.
IV-19
-------
Table IV-8
Incremental Cost of Gas Due to Pollution Control for the Synthane Process
Accounting
Method
DCF*
Utility Financing
1.6% Sulfur Coal
Direct Combustion of Char
Incremental
Annual Cost of
Gas($/yr)
14,444,900
11,135,500
Incremental
Cost of Gas
(S/106 Btu)
0.19(0.206)
0.15
Indirect Combustion of Char
Incremental
Annual Cost of
Gas($/yr)
19,356,000
14,517,300
Incremental
Cost of Gas
($/106 Btu)
0.25(0.271)
0.19
Accounting
Method
DCF*
Utility Financing
2.9% Sulfur Coal
Direct Combustion of Chart
Incremental
Annual Cost of
Gas($/yr)
16,633,600
12,968,900
Incremental
Cost of Gas •
(S/106 Btu)
0.22 (0.247)
0.17
Indirect Combustion of Char t
Incremental
Annual Cost of
Gas($/yr)
22,103,800
16,724,300
Incremental
Cost of Gas
($/106 Btu)
0.29 (0.320)
0.22
The incremental cost ol" gas without considering by-product credits is given in parentheses.
'The char fuels a furnace which raises steam to satisfy process steam and power requirements.
IV-20
-------
O 1.6% SULFUR FEEDSTOCK
A 2.9% SULFUR FEEDSTOCK
.05
66
68
INCREMENTAL CAPITAL INVESTMENT (S X 106)
i
P
O
w>
O
P
CO
-------
V. THE CO2 ACCEPTOR PROCESS
The CO2 Acceptor Process, an outgrowth of coal gasification
research performed during the 1950's by the Consolidation Coal Company,
is unique in its use of lime or calcined dolomite. In some high Btu
processes, air is introduced to support the combustion which supplies
the heat required for gasification. This requires a nitrogen barrier to
enable the product to attain the heating values required of pipeline
quality gas. The lime or dolomite reacts exothermally with carbon
dioxide to supply the heat required for the steam-carbon reaction. It
also absorbs ("accepts") CO2 from the gasifier gases, thus increasing
the direct formation of methane and negating the need for a subsequent
shift converter.
To minimize the rate of "acceptor" re circulation, gasification
temperature is limited to about 816 degrees to 857 decrees C (1500 de-
grees to 1575 degrees F). At these relatively low gasifier temperatures,
only very reactive feeds, such as lignite and subbituminous coals, are
commercially suitable as feedstocks.
A pilot plant, processing 36-metric tons per day (40 short tons/
day) of coal, costing $9 million, has been constructed in Rapid City, North
Dakota, by the Sterns-Roger Corporation. Test runs began in the
spring of 1971 to provide data for future commercial designs. .Even-
tually several Western lignite and subbituminous coals will be pro-
cessed and a number of dolomites and limestones will be tested for use
as CO 2 acceptors. At present, only the gasification section is in-
stalled. After the process is demonstrated, gas purification and
methanation stages will be added.
1. PROCESS DISPLAY
The schematic flow diagrams and energy balance for the CO2
Acceptor process is shown in Figure V-l. The CO? Acceptor process
has been simplified after publication of the documents used as source
data for this synopsis. The overall emissions, however, are expected
to be similar to those developed here. The composition and flow rates
of the gaseous, liquid, and solid streams are presented in Tables V-l
and V-2.
V-l
-------
Table V-1
Compositions of Gaseous Streams
CO
Stream No./
Component. Vol 7i
0-,
CH4
C:H6 ' '
©
©
17.34 7.17
0.34
CO '. 14.10 8.23
CO-. 5.5!
H: 44.54
N-. 0.23
NH3 ' 0.80
H2S 0.05
COS 8.2 ppmV
H;O
SO2
Total
ft'3/hrX 10'6
Ib-mole/hr
gm-mole/sec
17.09
100.00
31.48
83.049
(10,473)
4.21
50.54
0.38
0.04
3.1 ppmV
29.43
100.00
18.85.
49.723
(6,271)
©
--
--
--
--
--
::
100.00
100.00
3.92
10.345
(1,305)
©
--
--
..
100.00
--
100.00
16.95
44.731
(5,641)
©
7.1 7
8.23
4.21
50.54
0.38
0.04
3.1.ppmV
29.43
100.00
37.47
98.856
(12,467)
©
21.0
--
--
--
79.0
--
--
100.00
41.07
108.361
(13,666)
©
--
3,52
29.77
64.88
--
::
1.80
0.03
100.00
49.93
131,733
(16,613)
©
19.65
0.39
15.98
5.34
50.49
0.26
0
21.88
0.43
17.79
0.46
56.20
0.29
(To)
--
41.92
--
0.05 * ' (--
WO ppmW} 0.45
9.3ppmV| )--
7.84
--
100.00
27.76
73,257
(9,239)
2.95
100.00
24.94
65.809
(8,299)
-57.63
--
100.00
3.26
8.605
(1,085)
CM)
22.46
0.44
18.26
0.47
57.69
0.30
0.38
--
100.00
24.30
64,112
(8,085)
©
--
81.65
0.09
0.89
5.15
0.58
--
11.64
--
100.00
15.45
40,754
(5,140)
(n\
--
92.40
•-
0.10
1.00
5.83
0.66
--
::
0.01
•--
100.00
10.93
28,837
(3,637)
©
0.19
0.06
0.02
--
22.06
0.17
54.44
--
0.01
23.03
0.02
100.00
90.03
237,547
(29,957)
©
21.00
--
--
--
--
--
79.00
--
(\6)
•-
--
--
85.92
--
--
--
> 9.33
--
100.00
21.28
56,148
(7,081)
4.75
--
100.00
1.45
3,828
(407)
Comp. of Liquid Streams (mol '•'< )
E
Water
C02
NH3
H-.S
Naphthalene
Total
Ib-mole/hr
gm-mole/sec
86.41
6.81
6.78
1 1 ppm
trace
100.00
9,777
(1,233)
-------
FIGURE V-l
CO,, Acceptor Process Flow Sheet
LEGEND
tph ions (metric or short) per hour
{ ] sulfur flow rate in metric tons per hou
whr waste heat recovery
K.O. drum knockoul drum
ENERGY BALANCE
CARRIER
COAL
TOTAL ENERGY INPUT
PRODUCT GAS
AMMONIA
SULFUR
INERT GAS
COAL BURNED OURING
FEEDSTOCK DRYING
COOLING WATER
OTHER
TOTAL ENERGY OUTPUT
X106KCAL/HR
4431.5
4431.5
2623.0
27.6
11.4
140.9
460.0
419.3
749.3
4431.5
•Specific treatment and recovery processes were selected (or thit analysis.
The jytiem contained within the dotted tine was defined by the process
developer.
-------
Table V-2
Solids Analysis
Stream No. /Component
Coal/Char, wt %
H
C
N
O
s
H2O
Ash
Total
Dolomite, wt %
MgCO3CaCO3
MgOCaS
MgOCaCO3
MgOCaO
Total
Ib/hr Dolomite
(kg/hr Dolomite)
Ib/hr Char/Coal
(kg/hr Char/Coal)
A
2.82
41.70
0.69
13.32
0.59
33.67
7.21
100.00
2,488,000
(1,128,850)
A
4.01
65.37
1.10
17.17
0.90
--
11.45
100.00
1,417,300
(643,060)
A
0.920
75.152
0.805
1.850
0.688
--
20.585
100.000
788,250
(357,650)
A
0.550
66.012
0.190
1.045
0.763
31.440
100.000
516,200
(243,210)
A
'--
--
--
--
0.92
46.04
53.04
100.00
2,559,600
(1,161,340)
A
--
0.08
43.86
56.06
100.00
3,587,200
(1,627,590)
A
--
--
--
100.00
100.00
2,088,800
(947,730)
A
--
--
100.00
100.00
3,093,650
(1,403,650)
A
100.00
--
--
--
100.00
187,650
(85,140)
A
--
23.44
--
--
4.53
--
72.03
100.00
100.00
100.00
98,100
(44,510)
225,280
(102,210)
A
--.
. --
--
--
5.21
--
94.79
100.00
18,045
(8,187)
-------
(1) Bases for Analysis
The analysis shown is for a 0. 59 percentage sulfur (as
received wet basis) North Dakota Lignite coal whose moisture-
free composition is as follows:
COMPONENT ANALYSIS (DRY COAL)
Components We
Hydrogen 4.25
Carbon 62.87
Nitrogen 1.04
Oxygen 20.08
Sulfur 0.89
Ash 10.87
Total 100.00
This 33. 67 percent moisture coal has a heating value of
3927 kcal/kg (7068 Btu/lb). The plant is sized to consume
1129 m ton/hr (1244 short tons/hr) of lignite feed and produce
7. 43 x 106 m3/day (about 260 x 106 ft3/day) of an 8474 kcal/m3
(952.3 Btu/ft3) gas. The plant's coal-to-gas thermal efficiency
is determined to be 59. 2 percent.
(2) Layout and Symbols
The general direction of process flow is from the coal
supply on the left to the product gas on the extreme right of the
flow diagram, with residuals and by-products shown along the
bottom. The bold line indicates the flow of the primary gasifi-
cation process.
s
The figures contained within the triangles, circles, and
squares, respectively, refer to the solid, gaseous, and liquid
stream compositions depicted in Tables V-l and V-2. The
bracketed figures report the sulfur flow rate for the 0. 59 percent
sulfur feedstock. The overall energy balance, shown to the
right of the flow diagram, is calculated in Table V-3. That
V-5
-------
portion of the CC>2 Acceptor Process defined by the process
developer is contained within the dashed line in Figure V-l; the
remaining portions of the flow diagram are proposed for this
analysis.
The rhombic-shaped units represent intermediate products
(such as cooling water makeup).and sources (such as blowdown)
for which the distribution is not shown. Nonintegral pollution
cleanup processes are indicated by sloping rectangles. The
treatment and recovery processes which this symbol represents
are described in the process description and pollution control
portions of this chapter. Inverted trapezoids denote residual or
by-product storage.
2. PROCESS DESCRIPTION
The unit processes which define the CC>2 Acceptor Process are
described in this section, and the generation of pollutants and manu-
facture of by-products is discussed where appropriate. The discussion
of pollutant control and cleanup processes are covered in the following
section.
(1) Coal Preparation
The raw lignite delivered to the plant is first crushed to
less than 5/8 cm in hot-gas-swept impact mills. Its moisture
content is reduced from 34 to 16 percent in this step. Flash
dryers further dry the lignite until it is substantially free of
moisture. In the final preparation step, fluidized preheaters
heat the dry lignite to about 300 degrees C (570 degrees F),
driving off some additional moisture and CC>2.
Multistage drying of the lignite conserves total heat, and
the dried, heated feed reduces the devolatilizer heat requirement.
Lignite, and lignite fines which are generated in the crushing
step, are used as fuel in the feed preparation section.
V-6
-------
(2) De volatilization
The coal-to-gas conversion system employs three fluidized
bed reactors: a devolatilizer, gasifier, and regenerator. Pre-
heated lignite feed enters the devolatilizer through lock hoppers.
The devolatilizer operates at 815 degrees C (1500 degrees F)
and 20. 4 kg/cm^ (290 psia). Its bed is fluidized by gases evolved
in the gasifier and by steam. Most of the heat requirement is
supplied by hot calcined domolite flowing into the devolatilizer
from the regenerator (stream /7\ ). Makeup dolomite is also
injected into this vessel. This calcined dolomite removes CC>2
from the devolatilizer gas by the exothermic reaction,
MgO • CaO + CO2 —> MgO • CaCO3
(stream /5\ ). Spent acceptor and lignite char are separated as
they leave the devolatilizer. Spent dolomite returns to the
regenerator for calcination, and the devolatilized lignite is trans-
ferred by superheated steam to the gasifier vessel (stream /§V ).
(3) Gasification
Approximately 44 percent of the residual MAF (moisture
and ash-free) char is gasified in the gasifier at 857 degrees C
(1575 degrees F). The amount of carbon conversion is dependent
on the quantity of carbon required for acceptor regeneration,
discussed in section (4) below. The gasifier is fluidized by steam
(stream (4) ) and by recirculated gasifier effluent gases
(stream (5) ). The heat required for gasification is supplied by
hot calcined dolomite from the regeneration vessel (stream /o\.).
Spent dolomite (stream /&.) and residual char are separated in
the gasifier and withdrawn in separate streams for transfer to
the regenerator. The gas produced is returned to the devolatilizer.
(4) Regeneration
In the regenerator, residual char is burned at 1060 de-
grees C (1940 degrees F) to generate the heat necessary for the
recalcination of the dolomite. Regenerator off-gases carry
unburned char and ash to cyclones for separation from the gas.
The calcined dolomite is recycled to the devolatilizer and gasifier
V-7
\ • ,
-------
vessels. A dolomite makeup rate of approximately 2 percent is
required to maintain a satisfactory level of dolomite acceptor
activity and to replace losses. Sufficient sensible heat is con-
tained in the regenerator off-gases to generate and superheat the
steam required in the gasifier. Expansion of the cooled regen-
erator gases from 593 degrees C (1100 degrees F) and 19 kg/cm2
(270 psia) through a turbine to 1.1 kg/cm2 (15. 5 psia) and
237 degrees C (459 degrees F) generates enough power to drive
the main air compressor and the product gas compressor.
(5) - Gas Purification
The main components of the devolatilizer effluent gas are
methane, carbon monoxide, carbon dioxide, hydrogen, and
steam. Laboratory experiments have confirmed that as a result
of its 815 degree C operating temperature, no tar or hydrocarbon
liquids are contained in the devolatilizer off-gas. The hot efflu-
ent gas is first partially cooled in waste heat recovery boilers
which generate process steam. The gas is then further cooled
to about 110 degrees C (230 degrees F) in a water-scrubbing
tower. Hot potassium carbonate scrubbing removes about
90 percent of the CO2 and essentially all of the I^S from the
gas. Cleaned gas is cooled to near 38 degrees C (100 degrees F),
and condensed water is separated. Final traces of sulfur are
removed in iron sponge towers. The gas is methanated by any
suitable methanation process. The Consolidation Coal Company
has indicated interest in the U. S. Bureau of Mines' tube-wall
reactor in which the gas is methanated by contact with flame-
sprayed catalyst on heat-exchange tubes, and also in the Bureau
of Mines' hot gas recycle methanation system. The final product
gas, nearly sulfur and carbon monoxide free, has a heating value
of 8474 kcal/m3 (952. 3 Btu/ft3). It is compressed to 70. 3 kg/cm2
(1000 psia), dehydrated, and then delivered to the pipeline.
(6) Dolomite Preparation .
About 2040 metric tons/day (2250 short tons/day) of |
dolomite are added to the devolatilizer as makeup to maintain1
activity of the acceptor and to replace losses. Dolomites and
limestones from Ohio, Virginia, and South Dakota have
been evaluated in the laboratory. Specific sources of dolo-
mite have not been selected in this report because
V-8
-------
these will depend upon gasification plant location and the activity
of dolomites available near it. The dolomite will be crushed to
finer than 16 mesh at the mine before shipping to the coal gasifi-
cation plant. Conventional solids transfer equipment can trans-
port dolomite from stockpile to lock hopper for injection as
needed into the devolatilizer.
(7) Off sites
The regenerator air compressor is driven by expansion of
regenerator hot effluent gases through a power turbine. For
this plant size, two 2-stage, axial compressors will be required
with inter-coolers; total horsepower requirement will be about
225, 000 horsepower.
Enough power may be developed by the regenerator gas
power turbine to drive the 33, 000-horsepower product-gas
compressor. Steam is generated by the devolatilizer off-gas, by
the regenerator off-gas, and by interstage steam generators in
the methanation system. This amount of steam may be somewhat
greater than the plant steam requirements. Any excess steam
will be used to generate power.
(8) Pollutant Generation and By-Product Manufacture
The solids separated from the regenerator off-gas are
lignite ash and dolomite fines. These solids contain approximately
70 percent of the lignite feed sulfur as MgO • CaS. Before dis-
posal, the solids must be stabilized to prevent oxidation and
hydrolysis of this sulfide. The ash is treated in a fluidized bed
with off-gas from the hot potassium carbonate regenerator
(stream QQ) ). This gas is mainly steam and CO2» containing
less than 0. 5 percent sulfur compounds. Carbon dioxide and
water react with the MgO • CaS to form H2S and MgO • CaCO^.
According to the process developer, the product here will be
MgCOs • CaCOs- With excess MgO • CaO reacting, the CO2
content in the off-gas (stream @ ) will be reduced improving
the Glaus plant operation. The stabilized solids can be slurried j
with water and pumped to ash settling ponds. j
I
Gases from the ash-treatment reactor are enriched with
H2S to about a 9-percent concentration (Stream @ )„ Although '
this concentration is low, it is acceptable for Glaus plant feed.
V-9
-------
The Claus plant will be more expensive, and conversion to sulfur
will be poorer, than if a higher sulfur content feed gas were
available.
Excess water from the devolatilizer off-gas scrubber contains
dissolved CC>2, ammonia, and hydrogen sulfide. This water is
stripped by a Chevron wastewater treatment process. The process
produces both an H^S-rich gas stream, which is fed to the sulfur
recovery plant, and anhydrous ammonia, which is recovered for
sale. Most of the regenerator combustion gas is used as inert
gas for the transfer and drying of lignite in the feed-preparation
section. This gas is eventually vented to the atmosphere from
the feed preparation section flash dryer. Part of the combustion
gas from the regenerator is used to transport char from the
gasifier to the regenerator.
(9) Energy Balance
'The overall energy balance for the CC>2 Acceptor Process
is presented in Figure V-l and derived in Table V-3. From this
energy balance, the coal-to-gas process efficiency is found to
be
10, 408. 6 x 106 Btu/hr •=• 17, 585. 2 x 106 Btu/hr = 59. 2%
(10) Sulfur Balance
The sulfur balance for the 0. 59 percent sulfur* North
Dakota Lignite coal feedstock is quantified in Table V-4 and dis-
cussed in Section 3 of this process summary. The 6. 6 m tons/hr
of sulfur reported is the total amount evolved from all sources.
As shown in Figure V-l, the ultimate disposition of this
6.6m tons/hr of sulfur is 5.13 m tons/hr recovered as
elemental sulfur for sale, 0. 43 m tons/hr contained in the ash
returned to the mine for landfill, and 1. 1 m tons/hr released to
s
the atmosphere as SC>2.
Calculated on a dry, ash-free basis, the sulfur content is
1. 0 percent by weight.
V-10
-------
Table V-3
Energy Balance Calculations
Carrier
Calculations
1,825.5
1,663.8
2.973.5
Coal
Total energy input
2,488,000 Ib/hr x 7068 Btu/lb
Product gas
Ammonia
Sulfur
Inert gas
Coal burned during feedstock drying
Cooling water
Other*
Total energy output
P/hr x.:'
_ . JtonsTfi* x 2000 lb/fo1ft||g|KBtu/lb
5.66 tons/hr x 2000 Ib/ton x 3083 Btu/lb
49.93 x 106ft3/hr x 11.2 Btu/ft3
110,830 gal/min x 8.34 Ib/gal x 60 min/hr
x 30.0 Btu/lb
(by difference)
Includes heating values of other products, sensible heat of product streams, and heat lost to the atmosphere.
-------
Table V-4
Sulfur Balance
Carrier
Coal (lignite)
Total input
Sulfur (elemental)
Sulfur Dioxide to atmosphere
(0.64 short tons/hr in inert gas
+ 0.49 short tons/hr from
burning of fines)
Sulfides from Claus & Beavon
sulfur recovery
Sulfides from iron sponge drum
Sul fates in ash to minefill
Total output
Short Tqns/Hr.
' . -i7a*M •••••••
i 74
- ,.,;/: ipifv!:-
5,66 ;
1.13
0.06
0.02
0.47
7.34 ". V ;
Metric Tons/Hr.
6.66
6.66
5.13
1.03
0.05
0.02
0.43
6.66
3.
DISCUSSION OF POLLUTION CONTROL PROCESSES
The major waste streams of concern in the CO2 Acceptor
Process and their treatment are discussed in this section and sum-
marized in Table V-5.
(1) Sulfur Disposition
The path of sulfur, as it passes through this system, is
complex. The final disposition of sulfur in the dry lignite devola-
tilizer' feed, is:
Ten percent to devolatilizer off-gas as I^S to acid-
gas treatment
Ten percent to regenerator off-gas as SO0
&
Eighty percent to ash as MgO • CaS.
V- 12
-------
Table V-5
Source and Treatment of Major Waste Streams
Final Wastes
Source
Treatment
Lignite dust
Ash, waste dolomite
Wastewater
Hydrogen sulfide
Sulfur dioxide
Crushers, flash dryers, solids,
feed systems
Ash treatment
Scrubber, hot potassium
absorber, knockout drums
Ash treatment
Feed preparation, iron
sponge regeneration,
Beavon process
Cyclone separators,
bag filters
Immediate: ash ponds
Ultimate: worked out
strip mine
pits
Chevron process
Claus and Beavon
processes
None required - amounts
are below permitted
levels
The sulfur in the devolatilizer off-gas, together with
carbon dioxide, is extracted by hot carbonate purification. This
CC>2-rich gas (stream @) ) is used to stabilize the ash.. The
ash converts to MgO * CaCO^, and the combined sulfur from the
devolatilizer and ash (90 percent of the sulfur in the dried lignite)
goes to a Claus sulfur recovery unit as H2S.
The inert regenerator off-gas is used to purge the feed
preparation system. The sulfur in this off-gas, together with
the sulfur released in combustion of lignite fines in this system,
is exhausted to the atmosphere. A large portion of the sulfur
from the combusted lignite, however, is retained by the ash
because of its alkalinity. The total quantity of sulfur dioxide
released to the atmosphere is less than the EPA "New Source
Performance Standard" for coal-fired furnaces of
2.2kg SO2/106 kcal (1. 2 Ib SO2/106 Btu) when the total heat
release of the combustion and lime regeneratiop is considered.
The total SO2 emitted is 1. 323 kg/106 kcal (0. 735 lb/106 Btu)
of input lignite consumed.
V-13
-------
(2) Control of Solid Wastes
Lignite dust is generated during the crushing and subsequent
drying of the feed lignite. The bulk of these solids is separated
from the gas stream by two-stage cyclones. The fines collected
are burned to provide heat for lignite drying. The coarser solids
are returned to the system as feed. Two stages of cyclones
should be capable of removing about 98 percent of the total solids.
Approximately 56, 700 kg/hr (5 percent of the lignite feed) is
assumed to enter the first cyclone. Solids entering the bag filters
will equal about 1130 kg/hr. Bag filters will remove all but ,
about 0.1 percent of the entering solids. The quantity of solids
exiting from the bag filters to the atmosphere will amount to about
1.1 kg/hr (2.5 Ib/hr).
The coal ash contained in the spent dolomite is stabilized
in the ash treatment section by conversion of calcium sulfides to
calcium carbonate. The stabilized ash and dolomite are slurried
in water and pumped to retaining ponds for immediate storage.
Water will be recovered from the settled solids and recirculated.
When the pond is filled with solid material, it can be drained,
dried, and the solids excavated and transferred to underground
disposal in the worked-out strip mine pits.
Iron sponge guard drums can be regenerated several times;
the frequency of regeneration depends on the quantity of sulfur
contained in the feed gas. Eventually, the activity of the iron
sponge declines and it must be removed from the vessels. Spent
iron sponge should be regenerated before dumping to prevent
spontaneous combustion and to convert the iron sulfides to iron
oxides, which can then be discarded into the mine pits.
Methanation catalyst activity will decline as a result of
aging and occasional plant upset. The spent catalyst must then
be regenerated. The large amount of nickel contained in the
catalyst probably will permit economical return of spent catalyst
to the manufacturer for'nickel recovery.
V-14
-------
(3) Liquid Waste Streams
The stockpiled lignite at the plant should represent a
quantity equal to at least two weeks' supply of feed. Leachings
from this stockpile, as a result of rain and snowfall, can be
segregated and treated by biological oxidation.
The major process wastewater stream is excess water
from the devolatilizer off-gas scrubber (stream [T] ). Fortunately,
the devolatilizer does not generate oils or tars due to its oper-
ating temperature and, therefore, these need not be dealt with in
the wastewater treatment system. The scrubber wastewater
will be saturated with carbon dioxide, ammonia, and hydrogen
sulfide. The quantities present in the water, of course, are
dependent on the quantities in equilibrium in the feed gas stream.
The scrubber liquid may also contain some traces of naphthalene.
Ammonia, carbon dioxide, and hydrogen sulfide are separated
from the water by the Chevron wastewater treatment process.
The hydrogen sulfide is concentrated in the first stripping tower
and is sent to the Glaus plant for sulfur recovery. The ammonia
is stripped in a second stripping tower and is recovered as
anhydrous ammonia for sale. The wastewater may be treated
with active carbon for removal of traces of hydrocarbons or
phenols. The quality of the wastewater should be adequate for
the cooling tower makeup. The treated water may also be used
as makeup in the ash cooling and slurry transport system.
Cooling tower and boiler blowdown can be combined with
sanitary sewage and runoff water for conventional water treating.
A waste product from this treatment will be an inert solid resi-
due which can be discarded in the worked-out mine pit.
(4) Gaseous Waste Streams
Most of the gases vented to the atmosphere from the CC>2
Acceptor process come from the feed preparation and solids
feeding system, specifically the gases vented from the crushing
circuit, from the flash dryer exhaust, and from the lignite and
dolomite feed hoppers. The regenerator combustion gas com-
prises most of this gas; the remainder is from combustion of
lignite to preheat lignite dryer air. The regenerator off-gas
contains 0. 5 pounds of sulfur dioxide per million Btu
V-15
-------
r*
(0. 9 kg/10° kcal) of lignite char burned and, therefore, meets
the current emission standards. Lignite fines, having a sulfur
content of 0. 59 percent, are burned to preheat the air to the
dryer in the feed preparation section. The sulfur dioxide gen-
erated during this lignite combustion is 1. 07 pounds per million
Btu (1. 93 kg/106 kcal). The regenerator off-gas mixes with
lignite combustion gas in the dryer before being vented to the
atmosphere. The sulfur dioxide content of the combined streams
is 300 ppm, and averages 0. 735 pounds SO2/10^ Btu
(1. 323 kg/10^ kcal) released in the two process stages. Sulfur
dioxide removal is not required on this stream before it is
emitted to the atmosphere. The lignite ash retains a substantial
portion of the sulfur, thus reducing the sulfur emissions.
The iron sponge towers will be regenerated periodically
by oxidation with air. During this process, some SC>2 will evolve.
Because this operational step is intermittent, these exhaust gases
will be combined with the large flow of gas from the feed prepara-
tion and feeding systems for direct venting to the atmosphere.
Only a minor additional quantity of SO2 is evolved during the regen-
eration of the iron sponge drums (18 kg/hr of sulfur, 40 Ib/hr).
The gases vented from the hot potassium carbonate •'
regenerator contain less than 0. 5 percent H2S and about 42 per-
cent CC>2 (stream ^(5) ). This gas is utilized in the stabilization
of the calcium sulfide produced in the regenerator. The rejected
dolomite and lignite ash is treated by the CC^-rich gas in the
fluidized-bed ash treatment unit, converting the calcium sulfide
to calcium .carbonate and evolving H^S. The CC^-rich stream is
is thereby further enriched with H^S to slightly more than 9 per-
cent. One third of the H^S is oxidized in a Glaus boiler and the
SO2 formed reacts with the remaining I^S over bauxite catalyst
to form sulfur. The sulfur is condensed and sent to sulfur stor-
age. Tail gas from the Glaus plant is further treated in a tail
gas treatment process, such as the Ralph M. Parsons' Beavon
Process. The Beavon process further reduces the amount of
sulfur vented to the atmosphere. It is estimated that this com-
bined sulfur recovery system will convert 99 percent of the feed
sulfur to the elemental form, and only 1 percent will be vented
to the atmosphere.
V-16
-------
4. COSTS OF POLLUTION CONTROL *'
Costs required for control and treatment of pollutants evolved
during coal gasification by the CO2 Acceptor process are calculated
in this section. Since this process to produce pipeline gas is not yet
in commercial use, derived cost figures are estimates of costs for
similar pollution control unit processes. The major assumptions and
conventions adopted here are discussed in Chapter III. All by-product
flow rate data presented in Figure V-l also apply.
The incremental** capital investment for control of potential
pollutants and waste streams is presented in Table V-6. The invest-
ments considered are based on the control methods selected and
described in Sections 2 and 3. The incremental annual operating costs
for these emission control systems are presented in Table V-7. The
incremental capital investment using utility financing is $22. 6 million;
using the discounted cash flow (DCF) method, it is $24 million. The
annual operating costs, after taking credit for sulfur production at
$10 per long ton and ammonia production at $25 per short ton, is about
$1.5 million per year.
The derivation of the formulae for calculating the incremental
cost of gas production due to pollution control is presented in Chap-
ter III. For the DCF method, the required incremental annual cost of
gas, X, for the assumed rate of return is:
X = N + 0. 238161 + 0. 1275S + 0. 230777W,
where
N = incremental net operating cost
I = incremental plant investment
* The costs presented here are preliminary estimates; the actual
costs of emission control in this process will not be known until
a full-scale facility has been constructed and operated.
** Incremental costs in this report refer to that portion of the total
plant cost that is assignable to pollution control facilities.
V-l 7
-------
Table V-6
' ' ' £»
Incremental Capital Investment ($ Million)
Item
Incremental plant investment"
Dust emission control
Wastewater treatment
Sulfur recovery from ash
Claus sulfur recovery with tail gas treatment
Subtotal plant investment
Project contingencies0
Incremental plant investment (I)
Startup costs (S)d
Interest during construction6
Incremental working capital (W)'
Incremental Capital Investment (C)
$ Million
2.2
10.8
0.4
2.8
16.2
2.4
18.6
0.6
Utility
3.1
0.3
22.6
18.6
0.6
DCF
4.4 '
0.4
24.0
Notes:
a. Incremental plant investment, return on investment during con-
struction, and working capital are treated as capital costs in
year 0 (the year ending with completion of startup operation).
b. Installed costs, including engineering design costs, contractors'
profit, and overhead.
c. Includes costs for unexpected site preparation and hardware
requirements at 15% of plant investment.
d. At 20% of incremental gross operating cost.
e. For the DCF method, computed as the discount
rate x incremental plant investment for 1.875 veers'
average construction period. I( 1+i)n=I( 1.12)'-°7^= 1.236761 or
0.236761 additional investment.
For the utility financing method, computed at the
interest rate on debt x incremental plant investment
x 1.875 yrs.
f. Sum of materials and supplies at .9% of incremental
plant investment and net receivables at 1/24 of annual
incremental revenue.
V-18
-------
Table V-7
Incremental Annual Operating Cost ($)
Labor
Direct operating labor (3 men/shift x $5/hr
x 8304 shift hrs/man yr)
Maintenance labor (1.5% of I)
Supervisory (15% of direct operating and
maintenance labor)
Incremental labor cost
Administration and general overhead (60% of
incremental labor)
Other direct costs*
Supplies
Operating (30% of direct operating labor)
Maintenance (1.5% of I)
Incremental cost of supplies
Local taxes and insurance (2.7% of I)
Incremental gross operating cost
By-product credits
Sulfur (135.8 short tons/day x 365 days/yr
x 0.9 capacity factor x S10/LT
\2000 It/2240 short tons)
Ammonia (135.8 short 'tons/day x 365 days/
yr x 0.9 capacity factor x $25/short ton)
Less by-product credits
Incremental net operating cost (N)
124,600
279,000
60,500
37,400
279,000
464,100
278,500
1,500,000
316,400
502,200
398,400
1,115,600
3,061,200
-1,514,000
1,547,200
This item includes catalysts and chemicals expended and utilities purchased for the pollution
control processes.
V-19
-------
S = startup costs
W = incremental working capital.
For the utility financing case, it is:
X = N + 0. 1198C + 0. 0198W
where
C - incremental capital investment
N and W = as defined previously.
For this analysis, the annual gas production, G, is:
(10. 93 x 106 ft3/hr)(952. 3 Btu/ft3)(0. 9 capacity factor)(8760 hrs/yr) =
82. 062 x 1012 Btu/yr (20. 680 x 1012 kcal/yr).
This can be considered a typical production rate for a full-scale com-
mercial plant. Therefore, the incremental cost of gas, due to pollution
control is:
annual cost of gas _ X X
annual gas production G 82.062 x I0li! Btu/yr
Solution to the above equations for the incremental cost of gas
is shown in Table V-8. When using the DCF method, the incremental
cost of product gas, as a result of the investment and operating expenses
for waste system control, is 7. 5£/million Btu of gas produced. Using
the utility accounting method, it is 5. 2<*/million Btu.
Table V-8
Incremental Cost of Gas Due to
Pollution Control for the CO9 Acceptor Process
Accounting
Method
DCF
Utility financing
Incremental Cost of Gas
$/yr
6,151,300
4,260,600
$/106Btu
0.075
0.052
V-20
-------
5. REFERENCES
(1) "Pipeline Gas from Lignite Gasification, " R&D Report
No. 16, Interim Report No, 4, 9 May 1969, to OCR,
Department of the Interior.
(2) Fink, Carl A., "The CO2 Acceptor Process, " A paper
presented to the EPA Regional Conference, Denver,
Colorado, 1973 (unpublished).
V-21
-------
VI. THE LURGI PROCESS (PIPELINE QUALITY GAS)
The Lurgi gasifier, the only commercial gasification unit studied
in this program, was developed by the Lurgi Mineralbtechnik GmbH of
Frankfurt, West Germany. A unique feature of this process is that it
is the only fixed-bed type reactor system being considered for coal
gasification. Steam and oxygen are externally generated and supplied
to the gasifier. Steam is used to generate hydrogen while oxygen sup-
ports the combustion which supplies the heat to sustain the endothermic
gasifier reactions.
The present commercial Lurgi gasifier is limited to noncaking*
coals; this restriction may prohibit its use on many Eastern United
States coals. Recent gasifier research has been directed at mechan-
ical modifications to allow use of mildly caking coals. The commer-
cial application of such a Lurgi process may eventually become wide-
spread. Sixteen commercial Lurgi plants, producing a gas of about
3560 to 4000 kcal/m3 (400 to 450 Btu/ft3), have been built during the
last 30 years, and most are still in operation. The upgrading of this
gas to pipeline quality, though now being studied on pilot plant and
small commercial scale, has yet to be commercially proven. Meth-
anation problems, however, can probably be solved. The first of the
full-size plants, scheduled to produce 8. 16 x 106 m3/day (288xl06ft3/
day) of pipeline quality gas, has been designed for the El Paso Natural
Gas Company on a. site near Farmington, New Mexico. Process de-
sign work has been completed, a report on environmental factors has
been filed, and Federal Power Commission (FPC) authorization has
been requested to proceed with the project. Information contained in
this FPC filing was used in this chapter as an indication of one of the
process designs that could be employed with this gasifier.
In caking coals, volatile matter is liquefied at about 800 degrees
F (427 degrees C) causing the coal particles to agglomerate or
cake into large masses. Recent tests on the Lurgi facility in
Westfield, Scotland have indicated that a modified gasifier is
operable on mildly caking coals; however, a strongly caking
coal could only be processed at reduced capacity. At this
writing tests were not complete and therefore these results
are preliminary.
VI-1
-------
1. PROCESS DISPLAY
The schematic flow diagram and energy balance for the Lurgi
process is shown in Figure VI-1. Stream compositions, and flow
rates are presented in Table VI-1.
(1) Basis for Analysis
The process flow chart data, stream compositions, and
flow rate data presented in this process summary are based on
a Lurgi plant designed for the El Paso Natural Gas Company by
Lurgi and Stearns-Roger Incorporated and private communica-
tions with El Paso Natural Gas^' '3'. The process design de-
scribed in this synopsis has been simplified for presentation
and modified from the reference documents. Therefore this ,
presentation should not be regarded as the "El Paso" plant, but
rather as typical of what a Lurgi facility might accomplish.
The analysis shown is. for a 0. 69 percent sulfur (as re-
. ceived) coal from the Navajo Seam, whose composition is as
shown in Table VI-2. The heating value of this (16.25 percent
• moisture) coal is 4813 kcal/kg (8664 Btu/lb). Plant production
capacity is 8.16 x 106 m3/day (288 x 106 ft3/day) of an
8489 kcal/m3 (954 Btu/ft3) gas. The plant's coal-to-gas ther-
mal efficiency is determined to be 56.25 percent. Including
heat values of by-products, the overall thermal efficiency is
71. 6 percent.
(2) Layout and Symbols
The general direction of process flow is from the coal .,
supply on the left to the product gas on the extreme right of the
flow design, with residuals and by-products shown along the
bottom. The bold line indicates the flow of the primary gasifi-
cation process.
The encircled figures refer to the stream compositions
shown in Table VI-1. The overall energy balance is shown to
the right of the flow sheet.
VI-2
-------
FIGURE VI- 1
The Lurgi Process
Y
froeta Corxtonutt
LEGEND
I wan hitt botltr
ENERGY BALANCE
CARRIER
COAL
TOTAL INPUT
PRODUCT GAS
ASH
SULFUR
TAR
TAR OIL
PHENOLS
AMMONIA
NAPHTHA
COOLING WATER
LOSSES
TOTAL OUTPUT
ENERGY
-------
Table VI-1
Stream Compositions
STREAM COMPOSITIONS
STREAM NAMES
STREAM NUMBERS
Gas Components ( Vol. %]
O-,
CO;
H;S.COS,CS;>
C,H4
CO
H:
CH4
^1%
Ni + Ar
Ni
FLOW RATES do/hi)
Total Dry Gas
Water
Napnina
tar Oil
Tar
Crude Phenols
NH3
Coal(MAF)
Ash
Total Ib/ht
"(kg/hrr
Ratex 106ft3/day
Oxygen Gasifier
Coal Feed
1
314,950
1,250,3 10
373,220
1.938,480
• (879.530)
Air Gasifier
Coal Feed
2
---
67.533
268,053
80.012
415,587
(188,560)
Oxygen Gasifier
Ash
3
...
19.639
373,220
392.859
(178.250)
...
Air Gasifier
Ash
4
...
4,209
80.012
.84,221
(38.210)
Crude Gas
5
28.03
0.37
0.40
20.20
38.95
11.13
0.61
0.31
...
2.280,447
1 ,394.960
20.005
28.007
7,314
9.127
17,629
3.757.489
(1.704.850)
982.2
Gas Purification
Feed
6
32.36
0.34
0.39
11.70
43.63
10.70
0.59
0.29
2.450,001
2.680
20.005
2,472.686
(1. 121.910)
• 1079.7
Methanation
Feed
7
3.10
0.45
16.91
63.49
14.93
0.69
0.43
829.704
829.704
(376.450)
734.9
Methanation
Effluent
8
1.81
...
.01
4.16
92.93
1.09
513,694
1.316
...
. ...
:--:
' 515,010 '.
(233,670)
288.6
Pipeline
Gas
9
1.81
...
0.01
4.16
92.93
1.09
513.694
66'
513.760
(233.100)
288.6
N, Waste
to ATM
•10
0.86
99.14
1.577.467
9.995
...
1,587,462
(720,260)
511.0
Methanation
Condensate
II
...
22.58
77.42
« 69
314.625
314.694
(142.780)
Off Gas
to ATM
12A
2.29
85.45
SOppm
0.20
0.15
0.39
0.47
0.27
...
10.78
1,769,323
34,569
...
...
1.803,892
(818.460)
386.3
Off Gas to
Incineration
I2B
64.79
0.06
0.11
7.22
9.63
2.10
0.16
15.93
46,513
1,006
...
47.519
(21.560)
11.9
Combined
Acid Gas
13
96.12
1.06
0.22
0.41
0.76
0.59
0.30
0.54
1.670,209
10,900
...
1.681,009
(762,710)
35J.3
H2S From
Waste Liquor
14
95.91
4.09
8,853
8,870
...
17,723
(8.040)
1.8
Shift Effluent
15
...
36.95
0.32
0.35
5.03
46.80
9.75
0.53
0.27
1.392,164
357,765
10,939
15,314
3,999
4,991
9,640
...
1,794.812
(814,340
612.9
Shift Feed
16
28.03
0.37
0.40
20.20
38.95
11.13
0.61
0.31
...
1,246,945
762,764
10,939
15,314
3.999
4,991
9,640
2,054.592
(932.210)
537.1
-------
Table VI-2
Component Analysis (Dry Coal)
Components
Carbon
Hydrogen
Nitrogen
Sulfur
Oxygen
Trace compounds
Ash
Total
Weight %
58.73
4.30
1.02
0.82
12.12
0.02
22.99
100.00
Rhombic-shaped units represent intermediate products
(such as boiler feed water), uses (such as electric power), and
sources (such as process return water), for which the distri-
bution is not shown. Nonintegral pollutant cleanup processes
are indicated by sloping rectangles. The extensive treatment
and recovery processes which this symbol represents are dis-
cussed in the process description and pollution control sections
of this process summary. For a thorough in-depth explanation
of each licensed cleanup process mentioned, the reader is
referred to readily available process literature and the biblio-
graphy included in this report. Inverted trapezoids denote
residual storage.
2. PROCESS DESCRIPTION
The unit processes which comprise a possible Lurgi coal gasi-
fication system are described in this section. By-product manufac*
ture and pollutant generation are discussed where appropriate. The
discussion of pollution control processes is covered in the next section.
VI-5
-------
(1) Coal Preparation
Mined coal is crushed and screened to 3 x 0. 2 cm (approxi-
mately 1-1/4 x 1/8 inch) size. The sized coal is then conveyed
to coal storage bunkers atop pressurized coal lock hoppers.
Coal fines, generated during this preparation, are washed and
sold.
(2) Gasification
Coal is fed to the pressure gasifier (operating at a
pressure of 32. 3 kg/cm^ or 445 psig) through pressurized lock
hoppers. The oxygen and process steam are mixed and fed to
the bottom of the gasifier. The steam is the source of hydrogen
used in the process (H2O + C -» H2 + CO; H2O + CO -» H2 + CO2>.
Combustion of some of the coal with oxygen (C + O2 -^CO2)
supplies the heat required for gasification (621 to 760 degrees C
or 1150 to 1400 degrees F). The direct reaction of hydrogen and
coal produces some methane in the reactor (2H2 + C -»• CH^). As
the hydrogen-rich gas flows upward through this coal bed, the
coal is devolatized (the volatile matter is vaporized) and gasi-
fied to produce an 11.13 percent methane gas of the composition
shown in stream 5 of Figure VI-1*. This gas, which contains
tar, oil, naphtha, phenols, cyanides, ammonia,' sulfur compounds,
and coal fines proceeds downstream for further processing. De-
pressurization of the coal feed lock hoppers releases a fuel gas
which is collected in surge storage tanks, recompressed and
added to the primary gas stream. None is released to the atmo-
sphere. Ash is removed through ash-locks at the base of the
gasifier. This ash, 5. 0 percent by weight unreacted carbon
(stream 3) indicates 98.4 percent gasification of the original coal
feed as calculated by material balance. This carbon value is not
recovered but may be discarded with the ash. The ash is granu-
lar, nearly white, and may have by-product value as concrete
aggregate, or roadbed material. The sulfur content of the ash
should be negligible.
Source data did not quantify the composition of the hot, moist,
tar-laden gas effluent from the gasifier. Stream 5 represents
the composition of the cooled gas following quenching.
VI-6
-------
(3) Quench
The crude gas leaving the gasifier is quenched by a wash
cooler and passed through waste heat boilers to generate 8. 06
kg/cm2 (100 psig) process steam. Condensed gas liquor is
drawn off for further processing.
(4) Shift Conversion
Only about one-half of the total crude gas is subjected to
shift conversion (55 percent). Here, additional hydrogen is pro-
duced to adjust the H2:CO ratio for proper feed to the methana-
tion reaction. In the shift reactor, carbon monoxide is catalyt-
ically converted to carbon dioxide while steam is reduced to
hydrogen: CO + H2O t; CO2 + H2. In this process, the ratio of
the combined, shifted gas is 3. 7:1 because some CO2 is also
methanated. Some unreacted hydrogen remains in the product
stream too. Hot effluent gas from the exothermic shift reaction
is cooled in countercurrent heat exchangers with the feed gas.
The converted gas, together with the gas which bypasses the
shift converter, flows to the gas cooling area for waste heat re-
covery and cooling.
(5) Gas Cooling
Following shift conversion, the hot gas must be cooled to
facilitate subsequent purification. The cooling scheme used
here, as well as throughout the plant design, has been engi-
neered to minimize water consumption and to maximize the re-
covery and utilization of process heat by utilizing waste heat
boiler steam generators, boiler feed water heaters, aerial
coolers, and cooling water heat exchangers. This conservation
of heat energy results in a coal-to-gas plant efficiency of 56. 25
percent* (high compared to other Lurgi gasification plants). The
hot gas liquor and tar condensed from cooling the gas is sent to
separators where tar, tar-oil, phenol, and ammonia are re-
covered. The recovery of these impurities is discussed in
Section 3 of this process summary. The liquor from these
Derived in Energy Balance Section.
VI-7
-------
separators contains HCN which is generated during gasification.
This cyanide is withdrawn and sent to the sulfur recovery stage
where it is converted to thiocyanate and discarded.
(6) Purification
The acid-gas removal system, as indicated in the flow
sheet (Figure VI-1), is the Rectisol process. This process uses
methanol at low temperatures as the acid-gas absorbent. The
product from this process is a synthesis gas, free of impurities
which may otherwise harm the catalyst in the subsequent metha-
nation step. The Rectisol process is relatively complicated,
using a system of nine towers, and is not discussed in detail
here. The system design for this application was based on
nearly complete removal of COS from the product gas stream.
The gas is first chilled and then washed in methanol where
naphtha is removed. Next an absorber-regenerator set removes
a concentrated HgS stream (about 10 percent B^S) while the
majority of the CC>2 is removed in another final absorber-
regenerator pair. COS is dissolved in both absorbers. The
average of the two acid-gas streams is presented.as stream 13
in Figure VI-1 and contains 96 percent COo and 1. 06 percent
sulfides. Though not depicted here, actual plant design calls
for the splitting of these streams to maximize sulfur abatement.
The resulting CO2 and sulfides are delivered to the sulfur re-
covery plant. * The methanol is recovered, regenerated, and
recirculated, and the sulfur-free gas next flows to the metha-
nator.
(7) Methanation
The purified gas is reacted in a fixed bed reactor utilizing
nickel catalysts. Reaction heat from the exothermic methanation
* Both streams pass to separate Stretford absorbers for removal
of most of the H2S; COS and CS2 are not removed in the Stret-
ford facilities. The smaller, more concentrated stream (12B)
is incinerated; the larger stream (12A) is vented after the Stret-
ford with 80 ppmv sulfur. (It is further discussed in Section 3(1)
of this chapter.)
VI-8
-------
reaction is removed in a waste heat boiler and used to generate
process stream. Hot recycled gas is used to dilute the incoming
stream for temperature control. The output (methane rich) gas,
with a heating value of 8489 kcal/m (954 Btu/ft3) proceeds to
final gas cleanup.
(8) Compression and Dehydration
The methane rich gas is cooled and condensed water,
formed during methanation, is removed in a knockout pot. The
gas is then compressed by multi-stage centrifugal compressors
and cooled to 32 degrees C (90 degrees F) by air and cooling
water is meet pipeline gas specifications. A small quantity of
condensate is removed during compression and cooling; addi-
tional water is removed in a conventional glycol dehydrator.
These condensate streams are sent to the wastewater treatment
facility.
(9) Oxygen Supply
The high purity oxygen required in the Lurgi gasifier to
make a high-Btu synthetic gas, 5126 m tons/day (5650 short
tons/day), is manufactured on site using commercial air sepa-
ration processes. Water condensed in compressing and cooling
the air is recovered. A 99 percent N2 stream is purged to the
atmosphere (stream 10).
(10) Fuel Gas Production
About 17-1/2 percent of the 25, 626 metric tons/day
(28, 247 short tons/day) of coal feedstock supplied for the pro-
cess is used to produce a low Btu (1735 kcal/m , 195 Btu/ft )
fuel gas. This gas feeds the gas turbines and the boilers for
steam and electrical power generation. The gas is produced
in a gasifier similar to that required for the high-Btu crude gas
except that instead of oxygen, air is blown in to support the
combustion reaction and supply heat. The residue removed is
fine ash and some unreacted char (stream 4) which essentially
contains no pollutants.
VI-9
-------
(11) Steam and Power Generation
An on-site conventional power generating plant will supply
58.2 MW (including 8.7 MW exported to the mine) of continuous
electrical power in addition to the stream required for the plant.
This is the rated electric consumption at full capacity. The total
quantity of steam raised by the steam plant and process heat re-
covery units is 1, 902, 000 kg/hr (4,192, 010 Ib/hr). An important
feature of this facility is the use of gas turbines, fueled by a gasi-
fier product, to drive the electrical generators and plant com-
pressors.
Boilers combusting fuel gas and air generate superheated
steam which also provides motive power for turbine-driven com-
pressors as well as providing process reaction steam. The low-
Btu fuel gas, generated in the air-blown gasifier, is cleaned by a
high pressure (18.6 kg/cm , 250 psig) Stretford process (indi-
cated as Sulfur Recovery in Figure VI-1). Hydrogen sulfide re-
maining in the purified gas is less than 10 vppm, but the Stret-
ford process cannot significantly reduce the COS content in the
fuel gas. The composition of the treated fuel gas, fed to the
boilers, is shown in Table VI-3. The sulfur removed, prior to
combustion, results in power plant effluents (shown in Table VI-4)
whose SO2 level is 0. 2 kg/106 kcal (0.12 lb/106 Btu) of fuel
burned. This emission is a factor of 10 below the present per-
missible Federal level of 2. 2 kg/106 kcal (1. 2 lb/106 Btu) for
coal-fired plants. * Therefore, no further stack gas cleanup is
necessary.
This approach to power generation — gasification of coal,
followed by fuel gas desulfurization and combustion — results in
minimal sulfur emissions. However, it is penalized by high
capital requirements and greater coal consumption, when com-
pared to direct combustion of the coal.
The present Federal EPA specification of 1. 2 pounds/10 Btu
is applied to coal-fired boilers. However, the El Paso facility
was designed to comply with the more stringent New Mexico
emission regulations.
VI-10
-------
Table VI-3
Low^Btu Fuel Gas Composition
Components
C02
H2S + COS
C2H4
CO
H2
CH4
C2H6
' N2
Vol. %
14.86
.01*
.25
17.49
23.31
5.09
.38
38.61
*A small amount of naphtha present
in the gas is included. This naphtha !
, contains 0.21 percent sulfur.
Table VI-4
Composition and Quantity of Steam and Power
Generation Effluents to the Atmosphere
Components
Water vapor
Carbon dioxide
Nitrogen
Sulfur dioxide*
Oxygen
Nitrogen oxides (NO2)
Particulates
Ib/hr
270,047
645,384
5,095,317
334
1,020,800
562
Negligible
kg/hr
122,493
292,746
2,31 1,236
152
463,035
255
wt. %
3.8
9.2
72.5
14.5
334 Ib/hr of SO-, &
: ± ,- = 0.12 Ib SO2/106 Btu of fuel burned
2800 x 105 Btu/hr of fuel burned (0.216 kg SO2/106 kcal)
VI-11
-------
(12) Energy Balance
The overall energy balance for the process is presented
in Figure VI-1 and derived in Table VI-5. From this energy
balance the process efficiency, coal-to-gas, is found to be
11, 472 x 106 Btu/hr 20, 396 x 106 Btu/hr = 56. 25%.
Including the heat content of by-products raises the overall
energy recovered to 71. 6 percent.
(13) Sulfur Balance
The sulfur balance for the .0. 69 percent** sulfur Navajo coal
feedstock is quantified in Table VI-6 below and discussed in
Section 3 of this process summary. The reported sulfur output of
7369 kg/hr (16,246 Ib/hr) is the total amount of sulfur evolved from
all sources. As seen in Figure VI-1, the total sulfur output from
the process is composed of 76 kg/hr (167 Ib/hr) discharged to the
atmosphere from the Steam and Power Plant, 60 kg/hr
(133 Ib/hr) discharged to the atmosphere from the Stretford
sulfur recovery plant, 7068 kg/hr (15,582 Ib/hr) recovered for
sale as elemental sulfur, and 165 kg/hr (364 Ib/hr) contained
within the tar, tar oil, naphtha and ammonia by-products. Ex-
traction of the sulfur, remaining in these saleable by-products
by the seller is not considered. If further sulfur removal is re-
quired, it will be assumed to be performed by the purchaser.
(14) General Comments
The flow sheet reviewed here is an update of the flow
sheet presented in the report on environmental factors sub-
mitted by the El Paso Natural Gas Company . In this revi-
sion, the process has been simplified and the overall sulfur
emissions have been reduced, indicating the results of con-
tinued careful engineering by El Paso and Stearns-Roger, the
engineering contractor.
** Calculated on a dry, ash-free basis, the sulfur content is
1. 07 percent by weight.
VI-12
-------
Table VI-5
Energy Balance Calculations
Carrier
Coal
Total energy input
Calculations
8,664 Btu/lb x 28,249 ton/day (net after removal of fines)
x 2000 lb/ton-^24~hrs/day
Energy
(x 106 Btu/hr)
20,396
20,396
(x 106 kcal/hr)
5,140
5,140
Product gas
Carbon in ash (total)
Sulfur
Tar
Tar oil
Phenols
Ammonia
Naphtha
Cooling water
Losses*
Total energy output
954 Btu/ft3 x 288.6 x 106 ft3/day -^24 hrs/day
14086.8 Btu/lb x 23848 Ib/hr
3,983.4 BtuAb x 167 LT/day x 2,240 Ib/LT 4 24 hrs/day
16,970 Btu/lb x 88,824 Ib/hr
17,300 Btu/lb x 48,588 Ib/hr
14,021 Btu/lb x 11,271 Ib/hr
9,598 Btu/lb x 107,110 Ib/hr x 0.2
18,400 Btu/lb x 20,005 Ib/hr
62 x 106Btu/hr+ l,144x 106Btu/hr
(by difference)
11,472
336
62
1,507
841
158
206
368
1,206
4,240
2,891
85
16
380
212
40
52
93
304
1,067
,v
20,396
5,140
""Includes sensible heat of product, by-products and waste streams, air cooling, and other unaccounted losses.
-------
Table VI-6
Sulfur Balance
Carrier
Coal
Total input
By-products:
Tar
Tar oil
Naptha
Ammonia solution
Sulfur (elemental) ;
By-product total
Stretford sulfur plant effluents
discharged to atmosphere
H2S
«
CS2
SO2
COS
Turbine, boiler, and heater
effluents: 862 discharged
to atmosphere
Effluent total
Total output
Ib/hr
1 16,246
16,246
240
73
40
11
15,582
15,946
11
5
25
92
167
300
16,246
kg/hr
7,369
7,369
109
33
18
5
7,068
7,233
5
2
11
42
76
136
7,369 '
VI-14
-------
3.
DISCUSSION OF POLLUTION CONTROL PROCESSES
The major waste streams of concern in the Lurgi process and
their control treatment methods are discussed in this section and sum-
marized in Table VI-7.
Table VI-7
Nature and Treatment of Major Waste Streams
Final Wastes
Coal fines
Ash
Gas liquor
Hydrogen sulfide,
carbonyl sulfide,
and carbon disulfhtc
Sulfur dioxide
Amounts After
Treatment
259,722 kg/hr
77, 180 kg/hr
Tar, tar oil
phenols
ammonia
HCN
H2S
86 kg/hr
1 74 kg/hr
Sources
Coal feed, coat crushing,
screening, handling, lock
hoppers
Steam-oxygen and air-blown
gasificrs
Quench and gas cooling
Rectisol regenerator
Gas fired turbines, boilers,
heaters, incinerators
Treatment
Undersized coal
marketed
Sluiceway and
return to mine
for landfill
Separators, extractors,
Phenosolvan process
steam stripping
S tret ford process
None; amounts are well
below maximum permitted levels
(1) Sulfur Recovery
The H2S removed from the product gas during the gas
purification Rectisol process is treated in two low pressure
Stretford units to produce elemental sulfur. The fuel gas pro-
duced in the air blown Lurgi gasifier is similarly treated in a
high pressure Stretford unit. The Stretford process was de-
veloped by the British Gas Council and is now commercially
available through several U. S. licensors. The sulfur not re-
covered from the gas by the Stretford units appears either as
SO2 in the boiler stacks (shown in Table VI-4) or is vented to
the atmosphere directly (stream 12A) or after incineration.
The total sulfur compounds vented to the atmosphere by the
Stretford sulfur recovery units are identified and quantified be-
low in Table VI-8. These emissions are below the New Mexico
regulatory limits.
VI-15
-------
Table VI-8
Sulfur Compounds Vented to
Atmosphere from Stretford Units
Carbonyl sulfide
Carbon disulfide
Hydrogen sulfide
Sulfur dioxide
(after incineration)
Compound
(Flow Rate)
Ib/hr
172
6
12
50
kg/hr
78
3
5
23
Sulfur in Compound*
(Flow Rate)
Ib/hr
92
5
11
25
kg/hr
42
2
5
11
*Same as shown in Table VI-6.
Note: The Rectisol process sorbs approximately 80 percent of the carbonyl
sulfide into the CO2-rich gas. Only 20 percent of the original COS i
reports to the H2S rich gas which proceeds to incineration.
(2) Ash Disposal
Ash discharged from both the air and oxygen blown gasi-
fiers via ash locks is transported, with water discharged from
wastewater treatment, through a sluiceway to separation equip-
ment. Dewatered ash from the separation equipment is returned
to the mine for disposal and the slurry water is sent to evapo-
ration ponds (pond size not yet determined).
Alternatively, the clean ash from this system can be used
for concrete aggregate or roadbeds as is practiced by the South
African Coal, Oil and Gas Company (SASOL) and the Westfield
Lurgi plant in Westfield, Scotland.
(3) Gas Liquor Cleaning
The aqueous stream condensed during quenching and cool-
ing of the hot product gas and fuel gas passes through separators
where tar and tar oil is withdrawn. The coal tar (a dark, bitu-
minous, viscous liquid condensed and distilled from the gas
stream) has a heating value of about 8500 to 9500 kcal/kg
VI-16
-------
(15, 000 to 17, 000 Btu/lb). The tar oil (a lighter distillate
fraction containing cresote oil, anthracene oil and similar
aromatic compounds) has the same heating value as the coal
tar fraction. The clarified gas liquor is next treated in a
Phenosolvan process where organic solvents extract crude
phenols for sale. A deacidifier removes dissolved HCN, CC>2
and H2S and returns them for sulfur recovery (stream 14). The
HCN is converted to thiocyanate in the oxidizer section of the
Stretford system, and the thiocyanates are eliminated from the
system in the wastewater to evaporation ponds. This Stretford
bleed stream may also contain an undetermined amount of
thiosulfates and vanadates.
(4) Water Treatment
Raw water is subjected to chemical treatment by lime
softeners and then filtered. Low pressure boiler feedwater is
softened by zeolite treatment while water for process steam and
high pressure boilers is demineralized. The liquid effluent dis-
charges from water treatment are used as part of the carrier to
facilitate ash removal and are disposed of as land fill with the
ash or evaporated in evaporation ponds. The precise pollutant
composition of this contaminated liquid stream is not available
but it can be approximately quantified as:
818 1/min (216 gal/min) of lime treater sludge
2260 1/miri (597 gal/min) of boiler and cooling
tower blowdown
1245 1/min (329 gal/min) of contaminated gas
liquor
382 1/min (101 gal/min) of liquid waste from
plant's utility systems*
4705 1/min (1243 gal/min) Total
For example, plant firewater system and sanitary sewage
treating plant.
VI-17
-------
(5) Coal Fines
The Lurgi gasifier will not directly accept coal fines. The
feed to the gasifier is screened at 2 mm bottom-size and the
undersize material can be washed and marketed. The term
"coal fines" as used in this synopsis refers to undersize coal
too small for feed to the gasifier. The undersize coal in this
operation, however, may be acceptable feed for many other
process gasifiers. As much as 25-40 percent of the mine out-
put, depending upon mining and crushing techniques, will be too
fine for operation in the Lurgi gasifier. The sulfur in the waste
stream from coal washing will primarily be in the form of iron
pyrites. The ultimate disposal of pyrites might be accomplished
by blending them into the alkaline gasifier ash and then returning
the resultant mixture to the mine. This subject is discussed in
Chapter III of this report.
(6) Tars
The Lurgi gasification system generates significant quan-
tities of heavy tars. According to published data^', since these
tars contain only 0. 27 percent sulfur, they could be directed
combusted for heating in many locations. El Paso has investi-
gated the alternatives of either reinjection of the tar into the
gasifier for ultimate hydrocracking into oils and gas, or of par-
tial oxidation to produce synthesis gas in a manner similar to
that used in the Shell and Texaco Partial Oxidation processes
(discussed in Chapter X).
(7) Oils and Naphtha
The oils and naphtha by-products from the gasification
system (noted earlier in Table VI-6) contain small concentra-
tions of organic sulfides. This sulfur can readily be removed
from the by-products by distillate hydrodesulfurization. Hydro-
desulfurization techniques are discussed elsewhere in this re-
port.
VI-18
-------
4. COSTS OF POLLUTION CONTROL*
Costs of the pollution control and treatment processes described
in Section 3 are calculated here. The application of the Lurgi process
for high-Btu, sulfur-free gas is not a commercial reality. For this
reason the cost figures given in this section are based on:
Cost estimates provided by Lurgi Mineralbltechnik GmbH.
Lurgi, the process developer, has accumulated much cost
experience in low- and medium-Btu gas (200-450 Btu/ft3
or 1800-4000 kcal/m3) producing facilities which have been
commercialized for many years.
An economic evaluation, by Steams-Roger, Inc., for the
El Paso Natural Gas Company of a high-Btu Lurgi plant.
Costs derived from commercial experience for similar
pollution control processes.
The major assumptions and conventions adopted here are discussed
in Chapter III of this report. All by-product flow rate data presented
in Figure VI-1 apply. The value of tar, tar oil, and undersize coal
is included as part of the overall process cost rather than as by-
product credits from pollution abatement operations. These by-
products would be recovered in the basic operation of a facility even
if emission control were not practiced. In the alternative costing of
generating process power and steam is by direct combustion of coal,
however, the value of the reduced amount of tar and tar oil by-product
is considered as an incremental cost (at $0.50/10° Btu).
The incremental** capital and operating costs for pollution con-
trols are reported for the Lurgi process using indirect combustion of
coal (gasification of coal, and sulfur removal prior to combustion to
generate steam and electrical power needed in the process) and using
* The costs presented here are preliminary estimates; the actual
costs of emission control in this process will not be known until
a full-scale facility has been constructed and operated.
** Incremental costs in this report refer to that portion of the total
plant cost that is assignable to pollution control facilities.
VI-19
-------
direct combustion of coal. In the latter case, the sulfur effluent
appears as SC>2 and must be removed from the stack gases. For this
case it was assumed that the boiler house of a typical gasification
plant would comply with the Federal emission standards rather than
the more stringent New Mexico regulations.
Capital investment requirements are shown in Table VI-9 and
annual operating expenses are given in Table VI-10. Both capital and
operating costs of direct combustion of coal for steam and power
generation are less than those for indirect combustion (gasification)
of coal prior to combustion). The savings occur from not requiring
a second gasification system (the air-blown gasifier shown in Figure VI-1),
to supply the desulfurized fuel gas for steam and power generation,
costing in excess of $20 million, and from reduced maintenance and
labor requirements. Further savings may be realized by burning the
coal fires directly, rather than cleaning them for sale, if permitted
by local emission standards.
The derivation of the formulae for calculating the incremental
cost of gas production due to pollution control is presented in Chap-
ter III of this report. For the discounted cash flow (DCF) method,
the required incremental annual cost of gas, X, for the assumed rate
of return is:
X = N + 0. 23816 I + 0. 1275 S + 0. 230777W
where
N '= incremental net operating cost
I = incremental plant investment
S = startup costs
W = incremental working capital
For the utility financing case, it is:
X = N + 0. 1198C + 0. 0198W
where
C = required incremental capital investment
N = as defined above
W = as defined above
VI-20
-------
Table VI-9
—iumu.j-1 -r-iiJii-T -' T-im_i.._ Q
Incremental Capital Investment ($ Million)
Item
Incremental plant investment"
Coal fines cleaning
Phenosolvan phenol extraction
Gas liquor stripping (ammonia recovery)
Stretford sulfur recovery
Fuel gas production
Fuel gas cleaning
Incremental cost of steam and power
generation and compressed air
facilities
Limestone scrubbing for SC>2 recovery
Incremental cost of coal-fired boiler
Subtotal plant investment
Project contingencies0
Incremental plant investment (I)
Startup costsd (S)
Interest during construction6
Incremental working (capital' (W))
Incremental capital investment (C)
Indirect
Combustion of Coal
4.3
7.0
6.1
8.2
21.6
5.3
15.2
—
67.7
10.2
77.9 77.9
1.8 1.8
Utility DCF
13.1 18.4
1.4 1.8
94.2 99.9
Direct Combustion of Coal*
3.8
6.1
5.3
5.5 (high pressure Stretford
- - unit not required)
--
11.0
16.4
6.0
54.1
8.1
62.2 62.2
1.6 1.6
Utility DCF
10.5 14.7
1.0 1.3
75.3 79.8
Notes:
a. Incremental plant investment, return on investment during construction and working capital are treated
as capital cost in year 0 (the year ending with completion of startup operation).
b. Installed costs, including engineering design costs, contractors' profit, overhead, and licenses with no
contingencies.
c. Includes costs for unexpected site preparation and hardware requirements at 15% of plant investment.
d. AT 20% of incremental gross operating cost.
e. For the DCF method, computed as the discount rate x incremental plant investment for 1.875 years'
average construction period. I(l+i)" = 1(1.12) 1.875= 1.236761 or 0.236761 additional investment.
For the utility financing method, computed as the interest rate on debt x incremental plant investment
x 1.875 years
f. Sum of materials and supplies at .9% of incremental plant investment and net receivables at 1 /24 of
annual incremental revenue.
The costs for direct combustion of coal for power generation were derived in this study for comparison purposes.
VI-21
-------
Table VI-10
Incremental Annual Operating Cost ($)
Labor
I
to
Direct operating labor (I 2 men/shift
x S5/hr x 8.304 shift hrs./man yr.) "
Maintenance labor (1.5% of I)
Supervision (15% of direct operating and
maintenance labor)
Incremental labor cost
Administration and general overhead (60% of
incremental labor)
Other direct costs*
Supplies
Operating (30% of direct operating labor)
Maintenance (1.5% of I)
Incremental cost of supplies
Local taxes and insurance (2.7% of I)
Incremental gross operating cost
By-product credits (recovered in abatement process units)
Crude phenols (11,271 Ib/hr x S0.04/lb
x 0.9 capacity factor
x 365 days/yr x 24 hr/day)
Sulfur (167 LT/day x 365 days/yr x
0.9 capacity factor x S10/LT)
Ammonia (107.110 Ib/hr x 0.2,% NH
-=• 2000 short tons/lb x 24 hr/day x
365 days/yr x 0.9 capacity factor x 525/short ton)
Coal saved due to direct combustion (41,551 Ib/hr
at S227/ton)
Reduction in tar, tar-oil and naphtha recovered
@S0.50/10°Btu
Less by-product credits
Incremental net operating cost (N)
Indirect Combustion of Coal
498,200
1,168.500
250,000
1,916.700
1,150,000
2.500.000
149,500
1.168.500
1,318,000
2.103.300
8.988.000
3,554.400
548,600
2,111,100
-6,214,100
2,773,900
Direct Combustion Coal
(8 men/shift) 332,200
933,000
189,800
1,455,000
873,000
3,000.000
99,700
933,000
1,032,700
1,679,400
8,040,100
3,491,700
450,000 (137 LT/day)
2,073,800 (105,219 Ib/hr)
453,700
-195,600
-6,273,600
1,766,500
"This item includes incremental costs of chemicals, catalysts and make-up water.
-------
Annual gas production, G for the El Paso plant is:
(288.6 x 106 ft3/day) (954 Btu/ft3) (365 days/yr) (0.9 capacity
. factor assumed) * 90. 444 x 1012 Btu/yr (22. 792 x 1012 kcal/yr)
This can be regarded as a typical production rate for a commercial
Lurgi plant. Therefore, the incremental cost of gas, due to pollution
controls is:
annual cost of gas X X
12
annual gas production G 90.444 x 10 Btu/yr
Solutions to the above equations for the incremental cost of gas are
shown in Table VI-11. Though the direct combustion of coal is less
expensive, and uses 10 percent less coal, it releases 10 times as
much sulfur to the atmosphere. In the direct combustion of coal
after stack gas scrubbing, 1.2 Ib. 106 Btu (2. 16 kg/106 kcal) is re-
leased in the form of 803 the present permissible new source
standard whereas only 0.214 kg/106 kcal (0.12 lb/106 Btu) is
released from the Lurgi process.
Table VI-11
Incremental Cost of Gas
Accounting Method
DCF
Utility Financing
Direct Combustion of Coal*
Incremental Annual
Cost of Gas($/Vr)
17,077,100
10,807,200
Incremental Cost
ofGas($/106Btu)
0.19
0.12
Indirect Combustion of Coal t
Incremental Annual
CostofGas($/Yr)
21.971,500
14,086,800
Incremental Cost
ofGas($/106Btu)
0.24
0.16
"Combustion required for steam and power generator occurs at the gasification plant site.
tLow-Btu gasification of coal, followed by desulfurixation and combustion of gas.
VI-23
-------
REFERENCES
(1) Second Supplement to Application of El Paso Natural Gas
Company for a Certificate of Public Convenience and
Necessity, Proposed by Stearns-Roger Inc., FPC
Docket No. CP-73-131, October 8, 1973.
(2) "The Supply-Technical Advisory Task Force Synthetic
Gas-Coal, " Final Report, Federal Power Commission,
April 1973.
(3) Private communication. El Paso Natural Gas Company,
October 1973.
VI-24
-------
VII. THE HYGAS PROCESS
The Hygas process is an advanced coal gasification system which
can use all domestic coals including high volatile bituminous feeds.
The Institute of Gas Technology (IGT), in Chicago, started process
development of Hygas in 1946 under the sponsorship of the American
Gas Association (AGA). Largely due to this early development, Hygas
is currently the most highly developed of the advanced coal gasification
processes. From 1964 to the present, the U.S. Department of the
Interior's Office of Coal Research (OCR) has joined the AGA in sup-
porting this R&D effort. As of August 1972, $17 million had been
spent on process development.
The Hygas process, like most advanced coal gasification pro-
cesses, utilizes the hydrogasification reaction to produce methane.
Two-thirds of this product gas is produced directly in the prime
reaction. Maximizing direct methane formation improves the process
thermal efficiency and decreases the extent of catalytic methanation.
The gasification reactor operates at 70 to 105 kg/cm2 (1000 to 1500 psi).
Coal is gasified in two fluidized bed stages: a low-temperature stage,
with temperatures of 650 degrees to 760 degrees C (1200 degrees to
1400 degrees F), gasifies the most reactive portions of the feed and
forms large amounts of methane directly; the second stage, operating
at up to 1010 degrees C (1850 degrees F), gasifies the less reactive
portions. Residual carbon from this stage is withdrawn and used to
generate the hydrogen needed for hydrogasification. A mixture of
steam, hydrogen, carbon monoxide, methane, and other reaction
products serve as fluidizing gases. The steam also aids in controlling
gasifier reaction temperatures.
Three alternatives have been studied on ways to generate hydro-
gen from the carbon residue. These include:
Electrothermal gasification — heat required for the
steam-carbon reaction is supplied by passing an electric
current through the bed of fluidized solids
Steam-oxygen gasification — partial combustion of the
residual char with oxygen supplies the heat required for
the steam-carbon reaction
VII-1
-------
Steam-iron reaction—hot reduced iron reacts with
steam to form a hydrogen-rich gas. The iron is reduced
in a separate vessel by gas generated from char by par-
tial combustion with air and steam.
The steam-oxygen method was selected for this analysis since
it will probably be the first system commercialized. In the long term,
however, the steam-iron process holds promise for being the most
economical.
As with all high-pressure gasification systems, introducing feed
solids and removing ash solids from the pressurized gasifier require
process development. For this analysis, a slurry feed system is used
instead of a high pressure lock hopper system (which is not yet proven).
Solids are introduced into the gasifier as a slurry with oil; solids are
removed by depressurizing in a slurry with water.
In 1968, a detailed design was started for a $7 million Hygas
pilot plant to further process development. Construction of this pilot
plant was completed at IGT in early 1971 by Procon, Inc. This pilot
plant has now operated at design capacity producing pipeline quality
gas from 68 m tons/day (75 short tons/day) of coal. Tests to date
have used North Dakota Lignite feedstock. In September 1972, the
AGA/OCR awarded a 34-month contract to continue plant operation
with other ranks of coal and to install the steam-oxygen method of
hydrogen production in the pilot plant. Procon has also completed a
preliminary design of a demonstration plant to convert 4540 m ton/day
(5000 short tons/day)of bituminous coal to 2. 3 x 106m3/day (80 x
106ft3/day) of pipeline gas.
A full-scale commercial plant, producing 7. 1 x 10m
(250 x 106ft3/day) of pipeline gas, could be operational by 1981.
1. PROCESS DISPLAY
Schematic flow diagrams were developed for the Hygas process
using a Montana subbituminous coal feed (Figure VII-1) and an Illinois
No. 6 coal feed (Figure VII-2). Energy balances for each case are in-
cluded on the schematics and calculated in Section 2 of this analysis.
Stream compositions and flow rates are presented in tabular form fol-
lowing Figures VII-1 and VII-2. Design basis is a plant to produce
7. 1 x 106m3/day (250 x 106ft3/day) of pipeline gas.
VII-2
-------
FIGURE VII-1
The Hygas Process—Montana
Subbituminous
fUcvd*OUie.644l/mli
COAL TO PROCESS
COAL TO U-GAS BOILER
TOTAL INPUT
PRODUCT GAS
CRUDE PHENOL
AMMONIA
SULFUR
PRODUCT OIL
COOLING WATER EVAPORATION
OTHER LOSSES
TOTAL OUTPUT
•nd RKDWV PTOEWMI WOT Sriacttd by Smbx IGT Adviton lor Th«t Anriytii.
-------
FIGURE VII- 2
The Hygas Process—Illinois
No. 6 Coal
RKyd* Oil 16.525
LEGEND
CW cooling vwtcr
HP HOTM powtf
IMrin liitr/minut*
l_T long tool
BFW boiltr
-------
(1) Basis for Analysis
The process design contained in this report of the Hygas
process was specifically supplied for this study by the process
designer - IGT. This design is based on published and non-
proprietary data. The flow sheets depict a second-generation
Hygas plant which incorporates features that have been, or are
soon expected to be developed. Process flow designs were pre-
pared for both a low-sulfur western coal and a high-sulfur mid-
western coal. The moisture-free compositions of these selected
feedstocks are shown in Table VII-1. , ''
Table VII-1
Component Analysis (Dry Coal)
Component
Carbon
Hydrogen
Oxygen
Nitrogen
Sulfur
Ash
Total
Montana Subbituminous
(22% Moisture)
Weight, %
68.12
4.64
18.57
0.85
0.66
7.16
100.00
Illinois No. 6
(6.5% Moisture)
Weight, %
69.40
4.80
8.71
1.35
4.20
11.54
100.00
Product gas heating values for each feed are approximately 63 x
109kcal/day (250 x 109Btu/day). For the Montana subbituminous
coal, 21,481 m tons/day (23,678 short tons/day) of feed is re-
quired to produce an 8596kcal/m3 (966 Btu/ft3) gas. For Illinois
No. 6 coal, 15, 891 m tons/day (17, 517 short tons/day) of feed is
required to produce an 8569 kcal/m3 (963 Btu/ft3) gas. The coal-
to-gas thermal efficiency for these plants is determined to be about
60. 1 percent, and the overall efficiency to useful products is about
66-67 percent.
Higher thermal efficiencies have been reported in published
literature for this process. The figures quoted here have been
affected by the properties of the coal feedstock:
VII-5
-------
High oxygen content of the Montana coal consumes
some of the hydrogen that would have otherwise
been available for methane formation
High sulfur content of the Illinois coal also reacts
with some of the hydrogen thereby decreasing the
quantity of gas methanated
Oil-forming characteristics of both feed coals re-
strict the quantity of gas produced
Chemical characteristics of these coals for direct
methane formation during primary gasification are
lower than for some coals reported in the literature
U-Gas system for generating steam and power with
minimal emissions has inherent inefficiencies.
(2) Layout and Symbols
The general direction of process flow is from the coal
supply on the left to the product gas on the extreme right of the
flow diagram with residuals and by-products shown along the
bottom. A bold line indicates the flow of the primary gasifica-
tion process.
The figures contained within the circles refer to the solid,
liquid and gaseous streams described in Table VII-2 for the
Montana coal and in Table VII-3 for the Illinois coal. Numbers
are used to depict intermediate streams while letters denote
coal feeds, by-products, and residual products.
The Rhombic-shaped units represent intermediate products
and sources for which the distribution is not shown. Nonintegral
pollutant cleanup processes are indicated by sloping rectangles.
Inverted trapezoids denote residual or by-product storage.
2. PROCESS DESCRIPTION
The unit processes which define the Hygas process are described
in this section, with discussion of the generation of pollutants and the
VII-6
-------
Table VII-2
Montana Coal Stream Compositions
Stream name
Stream number
Temperature, °C
Pressure, kg/cm^
Component
CO
co:
H;
H:0
CH4
C2H6
C6H6
NH3
HCN
H2S
COS
N2
Phenol
Oil
Total Ib-mol/hr
(gm-mol/sec)
(x!06ft3/hr)
Raw Gas
1
316
81
Ib-mol/
hr
28.161
18.077
27,148
22.829
12,508
1,013
225
196
12
2,394
2.4
18
8
17,723
128.160
(16.162)
48.6
mol%
21.97
14.11
21.18
17.81
9.76
0.79
0;18
0.15
0.01
0.19
--
0.01
0.01
13.83
100.00
Oil Quench
Effluent
2
204
81
Ib-mol/
hr
28.161
18,077
27.148
22.829
12.508
1,013
225
196
12
2,394
2.4
18
8
2,246
112,683
(14.210)
42.7
mol %
24.99
16.04
24.09
20.27
11.10
0.90
0.20
0.17
0.01
0.21
0.02
0.01
1.99
100.00
Shift
By Pass
3
204
81
Ib-mol/
hr
9.387
6,025
9.049
7,620
4,169
338
75
65
4
79.8
0.8
6
3
748
37,570
(4.740)
14.2
mol%
24.99
16.04
24.09
20.27
11.10
0.90
0.70
0.17
0.01
0.21
0.02
0.01
1.99
100.00
CO Shift
Feed
4
--
81
Ib-mol/
hr
18,774
12,052
18.099
37,308
8.339
675
150
131
8
159.6
1.6
12
5
1,498
97,212
(12.260)
36.8
mol%
19.31
12.40
18.62
38.38
8.58
0.69
0.15
0.13
0.01
0.16
0.01
0.01
1.54
100.00
Combined
Shift Effluent
5
52
79
Ib-mol/
hr
13.177
33.069
42,140
29,928
12,508
1.013
225
204
4
239.2
2.6
18
.8
' 2,246
134,782
( 177000)
51.0
mol%
9.78
24.54
31.26
22.20
9.28
0.75
0.17
0.15
0.18
0.01
0.01
1.67
100.00
Scrub
Effluent
6
52
79
Ib-mol/
hr
13,172
32,721
42,124
180
12,501
1,013
218
--
228.2
2.6
18
6
102,185
(12,890)
mol %
12.89
32.02
41.22
0.18
12.24
0.99
0.21
0.22
--
0.02
0.01
100.00
Acid Gas
Effluent
7
38
70
Ib-mol/
hr
13,161
100
41,975
--
12,449
533
trace
18
68,236
(8,605)
25.9
mol%
19.29
0.15
61.51
18.24
0.78
0.1 ppm
0.03
100.00
Methanation
Feed
8
38
70
llwnol/
hr
13,161
100
41,975
--
12,449
533
--
18
68,236
(O05)
25.9
mol%
19.29
0.15
61.51
18.24
0.78
0.03
100.00
Product
Gas
9
38
70
Ib-mol/
hr
28
55
1,866
2
26,693
--
18
28,662
T3T6r5)
10.9
mol%
0.10
0.19
6.51
0.01
93.13
Note: All flow rates are rounded to the nearest mol/hr except for sulfur-bearing compounds. Volume flow rates are calculated at normal conditions of 14.7 psia, 60°F.
-------
Table VII-2
(Continued)
I
oo
Stream name
Stream number
Component
CO
co:
":
H-.O
CH4
C2«6
C6H6
NH3
HCN
H2S
COS
N2
Phenol
Oil
S02
Total Ib-mol/hr
(gm-mol/sec)
(x!06ft3/hr)
Foul Water
10
Ib-mol/
hr
5
348
16
43.625
7
7
204
4
8
8
44,232
(5,580)
0.013.
mol 7r
0.01
0.79
0.04
98.63
0.02
0.02
0.46
0.01
0.01
2ppmW
0.01
100.00
Oil
11
Ib-mol/
hr
3
2.240
2.243
(283)
4.8x1 0'3
mol%
0.13
99.87
I00;00
BTX-Oil
12
Ib-mol/
hr
218
1
6
225
(28)
4.8x 1 0"4
mol%
96.89
0.44
2.67
100.00
Acid Gas CO2
Vent to Stack
13
Ib-mol/
hr
11
32.089
149
52
480
0.3
1.3
32,783
(4,134)
12.4
mol %
0.03
97.88
0.45
0.16
1.46
5 ppmV
12ppmV
100.00
H2S Stream
to Claus Unit
14
Ib-mol/
hr
532
226.9
1.3
760
"(W
0.3
mol%
70.00
29.86
0.14
Scavanged
Gas
IS
Ib-mol/
hr
5
348
16
6
7
4
78
394
(50)
0.1
mol %
1.27
88.37
4.06
1.52
1.78
1.02
1.98
76 ppmV
100.00
Wastewater
16
Ib-mol/
hr
43,619
0.2
43,619
(5,500)
0.013
mol%
SppW
1 2 ppmW
lOppmW
20 ppmW
SO2 From
IT-Gas
17
Ib-mol/
• hr
257
67.3
324
(41)
0.1
mol%
79.25
20.75
100.00
Note: AH flow rates are rounded to the nearest mol/hr except for sulfur-bearing compounds. Volume flow rates are calculated at normal conditions of 14.7 psia, 60°F.
-------
Table VII-2
(Continued)
Stream name
Stream number
kg/hr
Component
C
H
O
N
S
Ash
Total
(gm-mol/sec)
Combined Coal
Feed
A
895.01 7 kg/hr
(22% Moisture)
wt%dry
68.12
4.64
18.57
0.85
0.66
7.16
100.00
Coal to
U-Gas System
B
175,388 kg/hr
(6.5% Moisture)
wt%dry
68.12
4.64
18.57
0.85
0.66
7.16
100.00
Coal to
Hygas System
C
57 1,258 kg/hr
(6.5% Moisture)
wt%dry
68.12
4.64
18.57
0.85
0.66
7.16
100.00
Ash
D
72,727 kg/hr
10. 2% Carbon
0.07% Sulfur
By product
Oil-BTX
E
32,154 kg/hr
Ib-mol/hr
C6H6 218
H2S 4
Oil 398
620
(78)
Crude Phenol
F
342 kg/hr
Ib-mol/hr
C6H6 -1
Phenol 8
15
(2)
Anhydrous
Ammonia
G
1,576 kg/hr
Ib-mol/hr
NH3 204
204
(26)
By product
Sulfur
H
4405 kg/hr
Ib-mol/hr
Sulfur 302.9
302.9
(38)
-------
Table VII-3
Illinois No. 6 Coal Stream Compositions
Stream name
Stream number
Temperature, °C
Pressure, kg/cm
Component
CO
CO2
H2
H2O
CH4
C2H6
C6H6
NH3
HCN
H2S
COS
N2
Phenol
Oil
Total Ib-mol/hr
(gm-mol/sec)
(xl06ft3/hr)
Raw Gas
1
. 316
81
lb-mol/
hr
23,217
17,359
24,167
19.501
15,260
502
149
628
39
1,361.6
11.8
59
8
14,765
17,027
(2.150)
44.3
mol %
19.84
14.83
20.66
16.66
13.04
0.43
0.13
0.54
0.03
1.16
0.01
0.05
--
12.62
100.00
Oil Quench
Effluent
2
204
81
lb-mol/
hr
23,217
17,359
24,167
19,501
15,260
502
149
628
39
1,361.6
11.8
59
8
2,079
104,341
(13,160)
39.5
mol%
22.25
16.64
23.16
18.69
14.63
0.48
0.14
0.60
0.04
1.30
0.01
0.06
0.01
1.99
100.00
Shift
By-Pass
3
81
lb-mol/
hr
7,739
5.786
8,056
6,500
5.087
167
50
209
13
453.9
3.9
20
3
693
34,781
(4.390)
13.2
mol %
22.25
16.64
23.16
18.69
14.63
0.48
0.14
0.60
0.04
1.30
0.01
0.06
0.01
1.99
100.00
CO Shift
Feed
4
52
79
lb-mol/
hr
15,478
11.573
16,111
35,920
10,173
335
99
419
26
907.7
7.9
39
5
1,386
92,480
( 1 1 .660)
35.1
mol "/,,
16.74
12.51
17.42
38.84
11.00
0.36
0.11
0.45
0.03
0.98
0.01
0.04
0.01
1.50
100.00
Combined Shift
Effluent
S
52
79
lb-mol/
lu-
ll, 304
29,294
36,106
30,459
15,260
502
149
654
13
1,357.9
15.5
59
8
2,079
127,260
(16.050)
48.2
mol'/'
8.88
23.02
28.37
23.93
11.99
0.39
0.12
0.51
0.01
1.07
0.01
0.05
0.01
1.63
100.00
Scrub
Effluent
6
52
79
lb-mol/
hr
11,299
28,946
36,090
165
15,253
502
142
--
--
1,348.4
15.5
59
6
93,826
(1 1.830)
35.6
mol %
12.04
30.85
38.46
0.18
16.26
0.54
0.15
1.44
0.02
0.06
0.01
100.00
Acid Gas
Effluent
7
38
76
lb-mol/
hr
11,290
88
35,971
--
15,212
119
trace
59
62,739
(7.910)
23.8
mol %
18.00
0.14
57.33
24.25
0.19
0.1 ppmV
0.09
100.00
Metha nation
Feed
8
38
76
lb-mol/
hr
11,290
88
35,971
15,212
119
59
62,739
(7.910)
23.8
mol %
18.00
0.14
57.33
24.25
0.19
0.09
100.00
Product
Gas
9
38
70
lb-mol/
hr
28
38
1,866
2
26,524
59
28,517
(3,600)
10.8
mol%
0.10
0.13
6.54
0.01
93.01
0.21
100.00
Note: All flow rates are rounded to the nearest mol/hr except for sulfur bearing compounds. Volume flow rates are calculated at normal conditions of 14.7 psia, 60°F.
-------
Table VII-3
( Continued)
-------
Table VII- 3
(Continued)
N>
Stream Name
Stream Number
kg/hr
Component
C
H
0
N
S
Ash
Total
gm-mol/sec
Combined Coal
Feed
A
662,139
(6.5% Moisture
wt%dry :
69.40
4.80
8.71
1.35
4.20
11.54
100.00
Coal to
U-Gas System
B
139,347 dry
wt% dry
69.40
4.80
8.71
1.35
4.20
11.54
Coal to
Hygas System
C
479,853 dry
wt%dry
69.40
4.80
8.71
1.35
4.20
11.54
Ash
D
84,974
10.37o of Carbon
0.3% of Sulfur
By-Product
Oil-BTX
E
18,047
Ib-mol/hr
C6H6 142
H2S 2.2
Oil 212
356.2
(45)
Crude
Phenol
F
581
Ib-mol/hr
C6H6 7
Phenol 7.8 .
14.8
(2)
Anhydrous
Ammonia
G
5050
Ib-mol/hr
NH3 653.8
653.8
(82)
By-Product
Sulfur
H
25,505
Ib-mol/hr
Sulfur 1753.8
1753.8
(221)
-------
manufacture of by-products given where appropriate. Pollutant con-
trol and cleanup processes are covered in the next section.
(1) Coal Preparation
Mined coal can be transported to the plant site by conveyor
belt, train, or truck. The coal stockpile will contain a 30- to
90-day plant feed supply. The coal is fed to crushers which re-
duce its size to finer than 8 mesh. A minimum amount of
100 mesh fines is produced. The coal is dried to near zero
moisture content in low-Btu gas-fired fluidized bed dryers.
Approximately 20 percent of the dried coal is utilized as feed
to the U-Gas system, which generates a low Btu fuel gas for
generating power, raising plant steam, and supplying miscel-
laneous plant fuel gas needs. This system is described in Sec-
tion 3 under "Steam and Power Generation. "
(2) Slurry Preparation
Dried coal is transferred to the slurry preparation section
where it is mixed with recycled oil from the gasification system.
Plunger pumps boost the coal/oil slurry to a pressure of 84 to
105 kg/cm2 (1200 to 1500 psia) to enable injection of the slurry
into the high pressure gasification system. Hot effluent gases
from the hydrogasifier fluidize a bed in the reactor where the
oil is vaporized. These hot gases supply the needed heat to
vaporize the slurry oil and dry the coal. The slurry dryer op-
erates at about 316 degrees C (600 degrees F).
(3) Gasification
The dried coal flows into the first of two hydrogasification
stages within the reactor. At 650 degrees C (1200 degrees F),
the coal particles react rapidly with the hydrogen-rich gas sup-
plied from the high-temperature gasification stage below
(stage 2). Large amounts of methane are formed by direct
hydrogenation of the more reactive components of the coal.
VII-13
-------
The method (now under development) of heating the coal and
contacting it with the reacting gases destroys its caking prop-
erties. This destruction of the agglomerating properties within
the gasifier reactor itself allows use of any feed coals without
requiring separate coal pretreatment. Coal from the first
Hygras reactor stage passes to the second stage, the high-
temperature gasification stage*
The high-temperature reactor operates at 954 degrees C
(1750 degrees F). Because the reactivity of the remaining coal
char is reduced, residence time required in this fluidized bed
is approximately 40 minutes. Heat required for gasification
in this stage is supplied by hot gases rising from the steam-
oxygen gasification reactor beneath and by exothermic heat of
the direct methanation reaction.
Char residue from the high-temperature reactor is trans-
ferred directly to a steam-oxygen fluidized-bed gasification
reactor. Hydrogen is formed by the reaction of steam with
carbon to form hydrogen and carbon monoxide. The heat re-
quirement for this endothermic reaction is supplied by injection
of oxygen which combusts part of the char. The method of in-
jecting the oxygen is designed to minimize burning of the hydro-
gen and carbon monoxide product gas and reduce the carbon
content of the ash to approximately 10 percent by weight. The
sulfur content of this ash is minimal. The high ash content
material is selectively removed from the steam-oxygen reactor,
slurried with water, and depressurized.
The ash-water slurry is transferred to settling ponds from
which clarified water is recovered and returned to the ash
slurrying system. Approximately 60 percent of the methane
contained in the pipeline gas product is formed within the hydro-
gas if ier.
(4) Oil Quench
As the first step in purifying the gasifier crude gas, it is
quenched by a cool recirculated oil stream to approximately
204 degrees C (400 degrees F). The final temperature of quench
is selected to avoid condensation of moisture. Most of the oil
is condensed, recovered, and returned to the slurry preparation
VII-14
-------
section for slurrying additional coal feed. Some heavy oil make-
up (from downstream units) is supplied to the oil quench section.
By avoiding condensation of water in the oil quench, emulsification
problems are avoided, cooling and reheating loads are decreased,
and the steam load to the following shift reaction is minimized.
(5) Shift Reaction
The next step in gas purification is to adjust the hydrogen/.
carbon monoxide ratio to about 3.1 in a shift converter. To attain
this ratio, about one-third of the feed gas bypasses this reaction.
The gas to be shifted mixes with steam at about 371 degrees C
(700 degrees F)« The mixture next passes into two stages of
fixed bed catalysts. A cob alt-molybdenum shift catalyst is used
here because it is tolerant of sulfur and oils in the feed gas, it is
rugged enough to withstand the high system pressure, and it can
withstand plant upsets during which hot catalyst might be inad-
vertantly contacted with water.
The shifted gas is cooled and water scrubbed at 52 degrees C
(125 degrees F) to remove most of its moisture content and essen-
tially all of the ammonia, phenols, and cyanide. Light oils and
BTX (benzene, toluene and Xylene) condensed with the water, are
separated and recovered.
(6) Acid-Gas Treatment
Cooled scrubbed gas enters a Rectisol solvent based acid
gas treating system. This system first removes the remaining
BTX, oils and water contained in the feed gas. Then, nearly
all of the acid-gas components are removed. The analysis
using Illinois No. 6 coal shows that CO2 is reduced from 30. 85
percent in the feed to 0.14 percent in the effluent. H2S is re-
duced from 1.44 percent to about 0.1 ppm. Even greater
cleanup is achievable when Montana coal is used.
Effluent gas from acid-gas treatment enters a fixed bed
sulfur guard chamber. This unit removes final traces of sulfur
from the gas before methanation. Because this gas contains
only a small amount of sulfur, the sulfur guard reagent remains
effective for long periods without requiring regeneration or
replacement.
VII-15
-------
(7) Methanation
Clean gas, now containing less than 0.1 ppm sulfur, flows
into the methanation system. The methanation feed gas contains
18 to 19 percent carbon monoxide which, when methanated, re-
leases large amounts of heat. The reaction of carbon monoxide
and hydrogen takes place in fixed bed reactors using a conven-
tional nickel-based methanation catalyst. This catalyst is
effective in the temperature range from 288 degrees C (550
degrees F) to a maximum of 482 degrees C (900 degrees F);
higher bed temperatures cause catalyst deactivation. The exo-
thermic reaction heat is controlled by use of cold, recycled
product gas. The exothermic heat generates steam for use in
the process.
(8) Dehydration
Following methanation, the gas is cooled, moisture is
condensed and the gas is dried in a conventional glycol drying
unit. After drying, the product eas has a heating value of
approximately 8587 kcal/m^ (965 Btu/.ft3y and is at pipeline
pressure of 70. 3 kg/cm^ (1000 psig).
(9) Water Supply and Waste Heat Recovery
Makeup water at the rate of 26, 500 1/min (7000 gal/min)
is brought into the plant from wells, surface streams, or reser-
voirs. Conventional water treating techniques are used to make
high purity boiler feedwater, auxiliary water, cooling tower
makeup water, and water for other plant needs.
Conventional induced draft cooling towers handle most of
the heat rejected by the process. Cooling towers receive
treated makeup water and cleaned treated wastewater from the
plant; 17,400 1/min (4600 gal/min) of water is evaporated in
the cooling towers, indicating extensive use of air cooling.
VII-16
-------
(10) Oxygen Supply
A conventional low-pressure air fractionation plant pro-
duces 98 percent purity oxygen for use in the steam-oxygen
gasification system. In the case of Illinois No. 6 coal feed,
2943 m tons/day (3244 short tons/day) of oxygen is needed;
for Montana subbituminous coal, 3292 m tons/day (3629 short
tons/day) is required. The oxygen plant prime movers are
driven by steam or gas turbines fueled by the U-Gas system.
(11) Energy Balance
The overall energy balance for the process is presented
in Figures VII-1 and VII-2. Table VII-4 summarizes the energy
balances for the two feedstocks selected for analysis.
(12) Sulfur Balance
The sulfur balance for both the 0. 51 percent* sulfur
Montana subbituminous coal and the 3.93 percent* sulfur
Illinois No. 6 coal are quantified in Table VII-5 and discussed
in Section 3 of this process summary.
For the case of Montana feed, the reported sulfur output
of 318. 6 Ib-moles/hour (40. 2 gm-moles/sec) is the total amount
evolved from all sources. This total is composed of: 307 Ib-
moles/hour (38. 7 gm-moles/second) of sulfur, 96.4 percent
recovery, either recovered as elemental sulfur or contained in
the saleable by-product oil (further removal of the sulfur in
this oil, if desired, is assumed to be performed by the pur-
chaser); 8.3 Ib-moles/hour (1.0 gm-moles/sec) as SO2 and COS
which is released to the atmosphere in the U-Gas Reactor stack
gas, the Rectisol acid-gas cleanup process and the Wellman-Lord
tail gas cleanup process; 3.1 Ib-moles/hour (0.4 gm-moles/sec)
of sulfur contained in the ash which ultimately goes to mihefill,
and 0.2 Ib-moles/hour (0.03 gm-moles/sec) of sulfur which is
contained in the wastewater.
When calculated on a moisture- and ash-free (MAF) basis, the
sulfur contents for the Montana and Illinois coals are 0. 71 and
4. 75 percent, respectively.
VII-17'
-------
Table VII-4
Energy Balance
Carrier
Input
Coal to process
Coal to U-Gas reactor
Total energy input
Output
Product gas
Crude phenol
Ammonia
Sulfur
Product oil
Cooling water evaporation
Other losses*
Total energy output
Montana Subbituminous Coal
x 106 kcal/hr
3,353
1,030
4,383
2,636
5
9
10
287
609
827
4,383
x 106 Btu/hr
13,307
4,087
17,393
10,461
20
34
39
1,138
2,415
3,286
17,393
Illinois No. 6 Coal
x 106 kcal/hr
3,360
975
4,335
2,620
5
27
56
162
603
862
4,335
x 106 Btu/hr
13,332
3,869
17,201
10,396
20
108
224
643
2,392
3,418'
17,201
•Includes sensible heat of product, by-product, and waste streams, as well as air cooling and other unaccounted losses.
Table VII-5
Sulfur Balance
(gm-moles/sec)
Carrier
Coal
Total input
By-products
Sulfur
In product oil-BTX
By-product total
Plant effluents : •
U-Gas stack gas to atmosphere
Wellman-Lord stack gas to atmosphere
Acid gas-regenerator CO? vent to
atmosphere*
In ash
In wastewater
Total output
Montana Subbituminous Coal
40.2
40.2
38.2
0.5
38.7
0.8
<0.1
0.2 (0. 1 6 gm-mol/sec as COS)
0.4
<0.1
40.2
Illinois No. 6 Coal
225.5
225.5
221.2
0.3
221.5
0.7(0.1 lbSO2/106Btu)
0.2 (250 ppm) (
1.0 (0.97 gm-mol/sec as COS)
2.1
<0.1
225.5
*IGT has conceived of a process to reduce these COS emissions by about 90%. This process is not presented in
this synopsis as patent protection has not yet been granted.
VII-18
-------
For the 1788 Ib-moles/hour (225.5 gm-moles/sec) of sul-
fur contained in the Illinois coal, though the recovery is greater
(98. 2 percent), the actual quantity of emissions is increased.
This is primarily due to greater sulfur in the ash from this feed,
and an increase in the release of COS from the acid-gas cleanup
system.
3. DISCUSSION OF POLLUTION CONTROL PROCESSES
The nature and treatment of the major waste streams generated
in the Hygas process are discussed in this section. Treatment methods
for controlling and recovering sulfur are also addressed. Waste
streams, their sources and treatment, are summarized in Table VII-6.
Table VII-6
Source and Treatment of Major
Waste Streams
Final Waste
Source
Treatment*
Coal fines
Ash
Wastewater
Hydrogen sulfide,
sulfur dioxide
CC<2 - rich gas
Nitrogen
Coal crushing, drying,
and transporting
Steam-oxygen gasifier:
U-Gas reactor
Water scrubber
Rectisol acid-gas treatment,
U-Gas cleaning, ammonia
recovery
Rectisol acid gas solvent
regeneration
Oxygen plant
Cyclone separators and
bag filters. Recycled
to process
Settling ponds and mine
disposal
Flash dryers, Phenosolvan
process, strippers, Phosam
process, two stage biological
treatment
Claus plant, incineration,
Wellman-Lord process
None required
None required
*For a more detailed explanation of the specific treatment processes discussed in this report, refer
to available published literature and to the reference section at the end of this report.
VII-19
-------
(1) Steam and Power Generation
For this process analysis, the method of motive and
electrical power generation and the raising of steam is considered
as one of the pollution control systems which limit the sulfur
released to the atmosphere. Twenty percent of the total coal
feed is gasified in the IGT U-Gas process where it is cleaned
of sulfur for use as a fuel to generate power, raise steam, and
other plant needs. The use of the U-Gas system significantly
reduces the sulfur content in the combustion gases.
The U-Gas reactor is a simple single stage, fluidized
gasifier utilizing steam and air to react with coal. The reactor
receives dried and sized coal from the coal preparation section.
This system operates at only 22.1 kg/cm^ (300 psig). This
pressure level enables the use of lock hoppers to feed reactor
solids. Gasification takes place at about 1038 degrees C (1900
degrees F). Higher localized temperatures are reached in the
fluidized bed by a submerged jet of steam and oxygen. Ash
from the U-Gas reactor is nearly depleted of carbon by the
high temperature gasification. The sulfur content is also
negligible. Ash is slurried with water, depressurized and
sent to ash disposal.
The high temperatures of the U-Gas reactor and method
of coal injection avoid tar and heavy oil production in the raw
effluent gases. The hot gases are cooled, thereby generating
steam, and then enter a proprietary hot gas treatment system
(the IGT Meissner process - included as part of U-Gas system)
which eliminates particulates and removes most of the sulfur
from the gas. This sulfur is regenerated as SC>2 and is directed
to the Glaus plant.
Part of the cleaned low-Btu gas generates 40 megawatts
of electrical power by combustion and expansion through a gas
turbine. Turbine exhaust gases raise steam in a waste heat
boiler, thereby generating additional power. This power is
used to satisfy in-plant needs. The gas turbines can supply
direct shaft horsepower to the oxygen plant. The remaining
low-Btu gas enters the plant fuel gas system. Most of this gas
is burned in the steam boiler plant which raises 170 kg/cm^
(2400 psig), 538 degrees C (1000 degrees F) steam, for use in
the gasification reactor and carbon monoxide shift section.
Some of the cleaned U-Gas product is burned in the coal prepa-
ration section to provide heat for drying the coal feed.
VII- 20
-------
Ash from the U-Gas reactor is slurried with water, de-
pressurized, and sent to ash settling ponds. Clarified water is
withdrawn from the ponds and recycled for additional ash
slurrying. When the ash ponds are filled, new ponds are put
into service and the filled ponds are drained and allowed to dry.
The settled ash is removed from the dried ponds and transferred
to mined out areas for final disposal.
(2) Sulfur Recovery
Rich solvent from the Rectisol acid-gas absorber enters
the regeneration section where some light oil and BTX is re-
covered. The regeneration section effectively separates dis-
solved H2S from CO2. The CO2~rich gas, vented to the stack
as stream 13, contains less than 0. 001 percent H2S and 0. 03
percent COS. The effluent from the lower sulfur Montana coal
contains only 0.004 percent COS. About 1.6 percent methane and
ethane are also present in the CO2 vent gas. This stream is vented
without further treatment. The H2S-rich stream, stream 14, con-
tains H2S in a concentration of 29.8 percent which is excellent for
Glaus plant feed.
Part of the H2S is oxidized in the Glaus plant boiler and is
combined with the remaining H^S stream and the SO2 stream
produced from the U-Gas acid-gas removal system. The com-
bined H2S-SO2 stream reacts over bauxite catalyst at near
atmospheric, pressure, to form elemental sulfur vapor. Two
reactor stages are used. The sulfur is condensed and sent to
storage as a liquid. When Illinois coal is the feedstock, 602
long tons/day (675 short tons/day) are produced; from Montana
coal 104 long tons/day (117 short tons/day) are produced.
Tail gas from the Glaus plant plus flash gases from foul
water are combined and incinerated. The sulfur components
oxidize to form SO2. The gases are then scrubbed with an
aqueous solution of sodium sulfite (the Wellman-Lord process
discussed further in Chapter III) to remove SO2. Scrubbed gas
will contain less than 250 pp'm of sulfur as SO2. The SO2 re-
covered from the sodium sulfite solution is recycled to the
Glaus plant for additional sulfur recovery.
VII-21
-------
Of the sulfur entering the plant with the Illinois coal,
about 98 percent is recovered in elemental form from the
Glaus plant, 1 percent is rejected in the ash, and about 1 per-
cent appears as stack gas effluent which is released to the
atmosphere from the steam boiler stack or the CC^-rich gas
vent.
Most of the sulfur is separated in the acid-gas treating
section of the Hygas plant or in the U-Gas cleaning system.
Smaller amounts are recovered from the foul water, and oil
streams. The 1 percent of the incoming sulfur discharged with
the ash is fixed into the lattice of the ash's remaining carbon.
With the Montana coal feed, due to the lower initial sulfur con-
tent, sulfur recovery is 95 percent and yet the concentration of
sulfur in the effluent streams is lower than with Illinois coal.
(3) By-Product Recovery
Some low-boiling oils and aromatics, such as benzene,
toluene, and xylene (BTX) are carried in the gas through the oil
quench section and are recovered in the water scrub and acid-
gas treatment sections of the plant (streams 11 and 12). The oil
is separated from water and sent to light oil storage. Heavier
oils are returned to the oil quench section to supplement the re-
cycle oil, and the lower boiling fraction is produced as by-product
oil. A BTX and oil production rate of 484,480 I/day (128,000 gal/
day) is predicted for the Illinois No. 6 coal, and 862,980 I/day
(228,000 gal/day) for the Montana subbituminous coal.
Water from the water scrubbing section contains a variety
of water soluble pollutants. These include phenol, dissolved
CO2 and H2S, and some hydrogen cyanide*. The foul water is
depressurized and degassed. Light oils are separated and the
foul water enters the phenol recovery section. Phenol is re-
moved by the Lurgi Phenosolvan process, a liquid-liquid
extraction process discussed further in Chapter III of this
report. Crude phenol is produced at approximately 15, 897
I/day (4200 gal/day). Water from the Phenosolvan process
is stripped of ammonia and H2S. This hot wet gas stream is
passed into a U. S. Steel Phosam process.
Most of the HCN manufactured in the gasifier is hydrogenated
to ammonia in the water-gas shift reactor.
VII-22
-------
In the Phosam process, ammonia is selectively extracted
from the gas by scrubbing with an aqueous ammonium phosphate
solution. Anhydrous ammonia is produced at 122 m tons/day
(134 short tons/day) from Illinois coal and 38 m tons/day (42
short tons/day) from Montana coal. After stripping, the waste-
water is expected to contain 5 ppm ammonia, 12 ppm hydrogen
cyanide, 10 ppm H2S, and 20 ppm of phenol and is of "satisfactory
purity for direct introduction into the plant cooling water circuit
(stream 16 ). The stripped gases flow to the Claus plant effluent
incinerator where cyanides and sulfides are oxidized.
(4) Control of Solid Wastes
Coal dust is generated during the coal breaking and
crushing operations. This dust is collected in two-stage
cyclones with bag filters to minimize solids lost to the atmo-
sphere. These fines are recycled to the system as feed
material.
As previously described, the coal ash obtained from the
U-Gas and the Hygas gasification reactors is sent to sealed ash
settling pounds as a slurry in water. When the ponds are filled,
the water is drained, and the ponds are dried before transferring
the ash to worked-out mining areas. Water is recycled or
.evaporated.
The sulfur guard chamber may contain metal-promoted
active carbon, or spent methanation catalyst which has been
rejuvenated for final sulfur removal. Only a tiny amount of
sulfur must be removed from the gas at this point, about 2. 3
kg/day (5 Ib/day) as I^S. Periodically, these catalysts will be
regenerated with steam and air. On regeneration of the active
carbon, the adsorbed sulfur compounds are converted to ele-
mental sulfur which is readsorbed. Regeneration of the spent
methanation catalyst evolves SO2 which is directed to the
Wellman-Lord system. Eventually, these solids must be re-
moved from the system and they should be oxidized before dis-
carding into the worked-out mining areas.
VII-23
-------
The methanation catalyst will require replacement occa-
sionally as the result of aging or deactivation due to misopera-
tion. This catalyst will probably be returned to the manufac-
turer for recovery of nickel and disposal of possible sulfur.
(5) Liquid Waste Streams
The stockpiled coal will cover a substantial area in the
plant. As a result of oxidation and rainfall, some organics and
sulfates will be leached from the pile and runoff water will
carry small amounts of coal fines. This water can be seg-
regated and the solids removed before biological treatment.
The main process water stream in this plant is the foul
water exiting from the ammonia recovery section of the plant.
This stream will be purified, as discussed earlier, and then
will feed the cooling water circuit. The cooling tower blow-
downs will be combined with other sanitary and process effluents
for conventional secondary treatment.
A maximum amount of treated wastewater will be recycled
to the process through careful plant design, utilizing low-tem-
perature heat sources within the plant, and innovative high
pressure steam boiler designs. Total recycle of this treated
wastewater may eventually be possible. Cooling tower and
boiler blowdown are combined with sanitary sewage and runoff
water for conventional water treating.
(6) Gaseous Waste Streams
The main gaseous waste streams vented to the atmosphere
are discharged from:
Coal preparation dryer vent
Steam boiler stack
Oxygen plant (nitrogen)
Acid-gas treatment facility (CC^)
Cooling tower (evaporated water)
Wellman-Lord treatment system (SO9).
VII-24
-------
The solids content in these gases is expected to be minimal.
The stack gases from the steam boiler will contain 0.2 kg/10°
kcal (0.1 Ib SC>2/million Btu) of fuel burned, significantly below '
the present Federal EPA nNew Source Performance Standard"
for coal fired facilities. The incorporation of the U-Gas system
into the coal-combustion scheme results in a significant reduction
in pollutants. The Wellman-Lord process stack gas will contain
about 250 ppm sulfur as SC>2. No potential pollution problems
are foreseen for the nitrogen rejected from the oxygen plant.
Water vapor from the cooling tower will normally be clean
steam. This could become troublesome only if leaks occur in
the process heat exchangers. Total combustible gases in the
acid-gas CC>2 vent equal about 2 percent by volume. This hydro-
carbon loss is a characteristic of the proved acid-gas removal
process employed in this preliminary design of the process.
The sulfur content of this stream may be as high as 300 ppmv,
primarily as COS, with the high sulfur Illinois coal. This sulfur
loss is based on expectations of the effectiveness of the acid-gas
process. At the cost of about 3 to 5 percent of the heating value
of the product gas, this CC^-rich gas stream could be incinerated
before venting which would convert the small COS content to SO2
and eliminate the hydrogen, methane, and ethane content. In
view of the present energy shortage, this incineration cost may
be too high.
4. COSTS OF POLLUTION CONTROL*
Costs required for control and treatment of pollutants evolved
during coal gasification by the Hygas process are calculated in this
section. Since this process to produce pipeline gas is not yet in
commercial use, derived cost figures are estimates of costs for
similar pollution control unit processes. The major assumptions and
conventions adopted here are discussed in Chapter III. All by-product
flow rate data presented in Figures VII-1 and VII-2 also apply.
* The costs presented here are preliminary estimates; the actual
costs of emission control in this process will not be known until
a full-scale facility has been constructed and operated.
VII-25
-------
The incremental* capital investment for control of potential
pollutants and waste streams is presented in Tables VII-7 and "VTI-8 '
for each of two feedstocks (a low and high sulfur content coal). The
investments considered are based on the control methods selected
and described in Sections 2 and 3. The incremental annual operating
costs for these emission control systems are presented in Tables VII-9
and VII-10.
The incremental capital investment for pollution control of the
low sulfur Montana subbituminous coal using utility financing is $57. 9
million; using the discounted cash flow (DCF) method, it is $61. 3
million. For the high sulfur Illinois No. 6 coal, the costs are $67.2
million and $71.7 million, respectively. The annual operating costs,
after taking credit for sulfur production at $10/long ton and ammonia
production at $25/short ton, is about $4. 8 million per year for the
Montana feed and $3 million per year for the Illinois coal. The
derivation of the formulae for calculating the incremental cost of gas
production due to pollution control is presented in Chapter III. For
the DCF method, the required incremental annual cost of gas, X,
for the assumed rate of return is:
X = N + 0238161 + 0.1275S + 0. 230777W
where:
N = incremental net operating cost
I = incremental plant investment
S = startup costs
W = incremental working capital.
For the utility financing case, it is:
X = N + 0.1198C + 0. 0198W
where:
C = incremental capital investment
N and W = same as defined above.
Incremental costs in this report refer to that portion of the total
plant cost that is assignable to pollution control facilities.
VII-26
-------
Table VII-7
Hygas Plant With Montana Subbituminous
Coal Feed, Incremental Capital Costs
Capital Requirements for Pollution Control3
U-Gas system
Dust control
Sulfur recovery unit
Sulfur recovery tail gas unit
Wastewater treatment
Subtotal plant investment
Contingency at 1 5%
Incremental plant investment (I)°
Startup costs (S)c
Interest during construction
Incremental working capital6
Incremental capital investment (C)
106$
25.4
2.2
0.9
2.3
10.8
41.6
6.2
47.8
1.1
Utility
8.1
0.9
57.9
47.8
1.1
DCF
11.3
1.1
61.3
Notes
a Incremental plant investment, return on investment during construction, and working
capital are treated as capital costs in year 0 (the year ending with completion of startup
operation).
b Installed costs, including engineering design costs, contractors' profit, and overhead, and
the contingency includes costs for unexpected site preparation and hardware requirements
at 15% of plant investment.
c. At 20% of incremental gross operating cost.
d. For the DCF method, computed as the discount rate x incremental plant investment for
1.875 years' average construction period. I(l+i)n = 1(1.12)1-875 = 1.23676 I or 0.23676 I
additional investment.
For the utility financing method, computed as the interest rate on debt x incremental
plant investment x 1.875 yrs.
e. Sum of materials and supplies at .9% of incremental plant investment and net receivables
at 1/24 of annual incremental revenue.
VII-27
-------
Table VII-8
Hygas—Illinois No. 6 Coal,
Incremental Capital Costs
Capital Requirements for Pollution Control3
106$
U-gas system
Dust control
Sulfur recovery unit
Tail gas recovery unit
Wastewater treatment
Subtotal plant investment
Contingency
Incremental plant investment I)"
Startup costs (S)c
Interest during construction (IDC)
Incremental working capital (W)e
Incremental capital investment (C)
25.4
2.2
4.1
5.8
10.8
48.3
7.2
55.5
1.3
Utility
9.4
1.0
67.2
55.5
1.3
DCF
13.1
1.2
71.1
Notes:
a. Incremental plant investment, return on investment during construction, and
working capital are treated as capital costs in year 0 (the year ending with
completion of startup operation).
b. Installed costs, including engineering design costs, contractors' profit, and
overhead, and the contingency includes costs for unexpected site preparation
and hardware requirements at 15% of plant investment.
c. At 20% of incremental gross operating cost.
d. For the DCF method, computed as the discount rate x incremental plant
investment for 1.875 years' average construction period.
I(l+i)n = 1(1.12)1:875 = 1.2367611 or 0.2367611 additional investment.
For the utility financing method, computed as the:interest rate on debt
x incremental plant investment x 1.875 yrs.
e. Sum of materials and supplied at .9% of incremental plant investment and
net receivables at 1 /24 of annual incremental revenue.
VII-28
-------
Table VII-9
Hygas—Montana Subbituminous Coal, '
Incremental Annual Operating Costs
Labor
Dollars
Direct operating labor (3 men/shift x
$5/hr x 8.304 shift hr/man-yr)
Maintenance labor (1.5% of I)
Supervision (15% of direct operating
and maintenance labor
Incremental labor cost
Administration and general overhead
(60% of incremental labor)
Other direct costs
Supplies
Operating (30% of Direct Operating Labor)
Maintenance (1.5% of I)
Incremental cost of supplies
Local taxes and insurance (2.7% of I)
Incremental gross operating cost
By-product credits
Sulfur (104 long tons/day x 365 days/year)
x 0.9 capacity factor x $10/long ton)
Ammonia (42 short tons/day x 365 days/year
x 0.9 capacity factor x $25/short ton)
Crude phenol
4200 gal/day x 7.5 Ib/gal x 365 day/year
x 0.9 capacity factor x $.02/lb
Total by-product credit
Incremental net operating cost (N)
124,600
717,000
126,200
967,800
37,400
717,000
754,400
341,600
344,900
207,000
893,500
967,800
580,700
2,088,700
754,400
1,290,600
5,682,200
(893,500)
4,788,700
VII-29
-------
Table VII-10
Hygas—I llinois No. 6 Coal,
Incremental Annual Operating Costs
Labor
Dollars
Direct operating labor (3 men/shift
x $5/hr x 8,304 shift hr/man-yr)
Maintenance labor (1.5% of I)
Supervision (15% of direct operating
and maintenance labor)
Incremental labor cost
Administration and general overhead
(60% of incremental labor)
Other direct costs
Supplies
Operating (30% of direct
operating labor)
Maintenance (1.5% of I)
Incremental cost of supplies
Local taxes and insurance (2.7% of I)
Incremental gross operating cost
By-product credits
Sulfur (602 long tons/day x
365 days/year x 0.9 capacity
factor x $10/longton)
Ammonia (134 short tons/day x
365 days/year x 0.9 capacity
factor x $25/short ton)
Crude phenol (4100 gal/day x
7.5 Ibs/gal x 365 days/year x
0.9 capacity factor x $0.02/lb)
Total by-product credit
Incremental net operating cost (N)
124,600
832,500
143,600
1,100,700
37,400
832,500
869,900
1,100,700
660,400
2,197,600
869,900
1,498,500
6,327,100
1,977,600
1,100,500
202,000
-3,280,100 -3,280,100
3,047,000
VII-30
-------
For Montana subbituminous coal, annual gas production, G, is esti-
mated to be:
(260 x 106 ft3/day) (966 Btu/ft3) (0. 9
capacity factor) (365 days/yr) = 82. 506 x
10*2 Btu/yr (20.792 x 1012kcal/yr)
For the Illinois No. 6 feed, annual gas production is:
(963 Btu/ft3) (259 x 106 ft3/day) (.9 capacity)
factor) (365 days/yr) = 81. 933 x 1012 Btu/yr
(20,647 x lO^kcal/yr).
This is regarded as a typical annual production rate for a commercial
Hygas plant. Therefore, the incremental cost of gas, due to pollution
control is:
Annual Cost of Gas _ X
Annual Gas Production G
Solutions to the above equations for the incremental cost of gas
are shown in Table VII-11. When using the DCF method, the incre-
mental cost of product gas, as a result of the investment and operating
expenses for waste system control, is 20^/million Btu of gas produced.
Using the utility accounting method, it is 14
-------
Table VII-11
Incremental Cost of Gas Due to
Pollution Control for the Hygas Process
Accounting
Method '
DCF
Utility financing
0.51% Sulfur Content Montana Subbituminus Coal
Incremental Annual Cost
of Gas, $/Yr.
16,567,000
11,742,900
Incremental Cost
of Gas, $/106Btu
0.201
0.142
Accounting
Method
DCF
Utility financing
3.93% Sulfur Content Illinois No. 6 Coal
Incremental Annual Cost
of Gas, $/Yr.
16,707,700
11,117,400
Incremental Cost
of Gas, $/106Btu
0.204
0.136
VII-32
-------
VIII. THE COED PROCESS
The Char-Oil-Energy-Development process (COED) was de-
signed to produce synthetic crude oil (syncrude) by pyrolysis of the
volatile components in bituminous coal. The by-products are a fuel
gas that can be purified for direct use, reformed to produce hydrogen,
or upgraded to pipeline gas; and a char that in some cases can be
used directly as a boiler fuel, or can be converted to a fuel gas.
Coal is crushed, dried, and heated to progressively higher tem-
peratures in a series of four fluidized bed reactors. The raw oil
evolved is purified and hydrotreated (treated with hydrogen) to remove
sulfur, nitrogen, and oxygen yielding a high grade synthetic crude
oil. An advantage of the COED process lies in its ability to process
agglomerating (caking) coals without prior oxidation that tends to re-
duce the yield of useful products.
The COED process has been under development since 1962 by
the FMC Corporation of Princeton, New Jersey, under the sponsor-
ship of the Office of Coal Research, U.S. Department of the Interior.
Following several years of development, and operation of a 45 kg/hr
(100 Ib/hr) Process Development Unit (PDU), the Blaw Knox Corpo-
ration of Pittsburgh designed and constructed a pilot plant in Princeton,
New Jersey. The capacity of the pyrolysis section of this pilot plant
is 33 metric tons (36 short tons) of coal per day; the hydrotreating
unit can process 30 barrels of oil daily. The successful operation of
the plant was demonstrated by a 30-day run in December 1970.
1. GENERAL DISCUSSION
All process flow data, flow rates, and stream compositions pre-
sented in this analysis of the COED process are based on the source
material referenced at the end of this summary and private communi-
cations with the process developers. * The analysis of the COED
* COED developers have suggested changes to the process as re-
ported here. Their comments are appended and referenced
where applicable.
VlII-1
-------
process is presented here for two feedstocks (Utah A Seam and Illi-
nois No. 6 coals) and two alternative methods of generating hydrogen
(steam reforming of pyrolysis off-gas and char partial oxidation).
The ultimate analysis of the specific feedstocks selected is reported
in Figures VIII-1 and VIII-2, along with other pertinent details. The
unit processes contained within the dashed lines were not defined by
the process developer but were assumed for this analysis. The sul-
fur and moisture contents, feed rates, and the respective yields are
given in Table VIII-1.
Table VIII-1
COED Feedstocks and Products'1
Carrier
Hydrogen generation
Feedstock
Sulfur content (dry %)
Moisture (%)
Feedrate (short tons/hr)
(m tons/hr)
Product
Char (short tons/hr)
(m tons/hr)
Oil (bbl/hr)
0/hr) '
Gas(X 103ft3/hr)
(m3/hr)
Utah A Seam
Gas reforming
0.5
6.0
443.7
402.5
206.2
187.1
591.7
94,063
820.8
23,245
Char oxidation
0.5
6.0
443.7
402.5
108.3
98.2
591.7
94,063
4,133.3.
117,055
Illinois No. 6
Gas reforming
3.7
11.0
462.5
419.6
230.2
208.8
377.9
60,075
562.5
15,930
*Other by-products include sulfur and ammonia. (See reference 5.)
2. PROCESS DESCRIPTION
The flow diagram and tabulations of associated material and
energy balances for the COED coal liquefaction process are presented
in Figures VIII-1 and VIII-2 for the case of steam-reforming to make
process hydrogen. The objective and operation of each step, including
emission generation, are described below under the following headings:
VIII-2
-------
FIGURE VIII-1
The COED Process—Utah A Seam
Coal
COMPOSITION OF SOLID AND LIQUID STREAMS IWT. %}
FEEDSTOCK
COMPONENT
C
H
O
N
S
ASH
TOTAL
MOISTURE
(nTIonZr
UTAH A SEAM
COAL
V
74,8
5.5
11.7
1.4
0.5
6.1
100.0
6.0
(402.6)
CHAR
85.2
1.1^,
1.0
1.4
0.7
10.6
100.0
-
(187.1 1
RAW OIL
A
83.3
8.7
6.5
1.0
0.3
0.2
100.0 *
-
(88.31
SYNCR
A
85.8
11.8
2.0
0,32
0.02
-
-
_
94.14
(85.4)
.
1
: /^
, <^UM «y-
| CHAR--^
,fFOR ILLINOIS COAL), •
COAL OuST ~~"j
J CO\TRQL
CO
f- _
74.9 tph
/ATMOSPHERE
| I
L-: —
i
187.1 tpri ;
\ CH4R / \ RESIDUE / \
P
^— — — _
CHEVRON^
TREATMENT* [
1
0.87 tph
AMMONIA /
-
\
)
STRETFORD 1
SULFUR
RECOVERY -
SuKidei
0,004 iph
Suilur
0.55 tph
SULFUR /
S02
ATMOSPHERE \
1
A
ZINC
OXIDE
v J
3-
^
21 kg/cm2
300 ps
" STEA
<9
M
REFORMING
,
1
SULFINOL
RE MO-
^ /
A
CO
CONVER-
SION
^ ;
Hydrogen
1
"- -1 "1
tph - M
1
om/hour
'Specific treatment and recovery processes ne>e defined by the process developer.
COMPOSITION OF GASEOUS STREAMS (VOL. "='
COMPONENT
H2
H20
CO
co2
C2H4
C H
2 6
C3H6
C3H8
C4HIO
°5
~2
""?
FLOW RATE
,03 A,
©@®©®®©®©
_ 37.58 48.96 41.93 95.15 1.30 49.30
15,50 16.29 18.18 13.08 2.08 ' 2.99 - 6.01 -
- 17.36 29.37 19.36 - 0.77 2267
13,98 13.85 13.56 12.94 14.69 14.44 _ 89.32 TRACE
14.14 4.90 15.78 4.85 0.34 21.06
- 0.35 0.38 - - 0,45
- 1.31 - 1.46 _ - 1.72
- 0.26 - 0.29 - - 0.34
- - - 0.17 - 0.19 - - 0.40
--- -- -- - 0.22
„__ ____ ,_
70.52 6986 68.24 - - - _ _
- - - 015 - 0.17 _ 1.88
62 ISO 222 210 130 139 55 23 23
103 ft3/hr j 2.200 5,300 7.800 7,400 4.600 4,900 1,930 BOO 820
-------
FIGURE VIII-2
The COED Process—Illinois
No. 6 Coal
"1
G
;
/
190°C
375°F
^
PYROL
STAC
315
600
^
540°C +
1000°F
^
>0°F
OLYSIS
ES M-llt
to°c
Char
PYROLYSIS
STAGE IV
870°C
1600°F
' • I M
74.9 tpt>
/ATMOSPHERE
Noie: values for Flow Rates and or Compositions refer to Illinois No. 6 Coal.
'Specific treatment and recovery processes ware defined bv the process developer.
COMPOSITION OF SOLID AND LIQUID STREAMS (WT. %) 1
FEEDSTOCK
COMPONENT
C
H
0
N
S
ASH
TOTAL
MOISTURE
short tom/ftr
(m loo/hr I
ILLINOIS NO. 6 i
COAL
\2/
67.0
4.7
11.0
1.2
3.7
12.4
100.0
11.0
462.50
1419.71
CHAR
V
73.4
1.3
0.3
1.2
3.1
20.7
100.0
-
230.20
1208.9)
RAW OIL
/6\
81.3
7.1
81
1.0
2.4
0.1
100.0
-
75.60
(68.61
SYNCR
S7.60
9.94
1.22
0.82
0.22
-
100.0
-
65.71
(5931
COMPOSITION OF GASEOUS STREAMS
VOL.%1
COMPONENT
H2
H20
CO
co2
CH<
C2H4
C2H6
C3H6
C3H8
C4H8
C4H10
C5
°2
HjS
FLOW RATE
103m3/h>
103 H3/Mr
© ®
38.5 1.3
-
18.4 0.7
13.0 75.3
12.3 0.4
0.6
10.1 0.3
1.6
0.2
-
1.5
-
_ _
3.8 22.0
98 17
3450 595
©
46.3
-
22.1
TRACE
,4.8
0.7 ;
12.2
1.9
0.2
-
1.8
-
-
TRACE
16
562
-------
Coal preparation
Pyrolysis
Oil recovery and filtration
Oil hydrotreatment
Gas purification
Hydrogen generation
Steam and power generation
Energy balance
Sulfur balance.
The corresponding pollution control practices and their costs are
covered in subsequent sections.
(1) Coal Preparation
Coal (stream \I7 and\&7 ) is ground by hammer mills
to minus 3 mm (1/8 in) size and conveyed to storage bins.
From the storage bins, the coal is fed to a dryer which is
heated by hot process gas (drying gas). The bulk of the coal
dryer exit gases is transferred to the first pyrolysis stage
and then to gas cleanup, eventually returning to the coal dryer
as fluidizing gas. Dried coal at 190 degrees C (375 degrees F)
is conveyed from the dryer to the first pyrolysis stage. Coal
preparation generates both coal dust and gaseous emissions.
These are discussed in Section 3 of the chapter.
(2) Pyrolysis
The pyrolysis operation is performed by heating coal
to progressively higher temperatures in a series of four fluid-
ized bed reactors. Pressure is maintained throughout at 1. 4
to 1. 7 kg/cm2 (6 to 10 psig).
In the first stage, the coal is rapidly heated to 315 degrees C
(600 degrees F) and devolatilized by fluidizing gas (stream (2) )
at 540 degrees C (1000 degrees F). The exit gases (stream (3) ),
containing some char and oil proceed to gas cleanup. Here the
gas is quenched with recycle water (an aqueous stream from gas
quenching in the oil recovery section). The oil-water mixture
is transferred from gas cleanup to the oil recovery section
where the oil is separated. The purified gas from gas clean-
up is recycled to the coal dryer. Another portion of this gas
VTTT- S
-------
is used to cool the product char from stage IV to 93 degrees C
(200 degrees F) before being vented to the stacks. Char, at
315 degrees C (600 degrees F), is next conveyed to the second
pyrolysis stage.
Pyrolysis stages two and three are contained in two sec-
tions of a vessel maintained at 455 degrees C (850 degrees F)
and 540 degrees C (1000 degrees F), respectively. The effluent
stream from the second stage (stream (4) ) contains the volatile
matter driven from the coal in the last three stages of pyrolysis.
This stream next flows to the oil recovery section.
Char from the third pyrolysis stage, at 540 degrees C
(1000 degrees F), is conveyed to the fourth stage, where it is
heated to 870 degrees C (1600 degrees F). The heat requirements
for the last three stages of pyrolysis (stages II, III, and IV) are
supplied by combustion of a portion of this char with oxygen
within stage IV. Based on the analysis of stage IV off-gas, steam
may also be added at this point (ref. 5). The remaining product
char from this stage (stream ^p and ^p ) is cooled in a fluidized
bed char cooler to 93 degrees C (200 degrees F) and sent to
storage.
(3) Oil Recovery and Filtration
Within the oil recovery and filtration section, the volatile
products from the second pyrolysis stage are first cooled from
455 degrees 'C (850 degrees F) to 77 degrees C (170 degrees F)
in a venturi scrubber. Next, most of the oil is condensed and
removed in a gas-liquid separator. The gas is then passed
through an electrostatic precipitator to remove the oil mist,
and finally through a spray tower to remove any remaining
traces of oil. The oil separated from the oil-water mixture
obtained from gas-liquid separation, the spray tower, and the
first stage pyrolysis gas cleaning unit is removed in a decanter.
The water is recycled to the gas-liquid separator and gas cleanup
of the first stage pyrolysis. The oil is dehydrated by heating
with steam coils and filtered mechanically to remove fine char
particles which escape the cyclone separators associated with
the pyrolysis reactors.
VIII-6
-------
(4) Oil Hydrotreatment
Treatment of the raw pyrolysis oil with hydrogen is re-
quired to remove unwanted nitrogen, sulfur, and oxygen and to
upgrade its specific gravity and viscosity to the level of about
25 degrees API syncrude. The raw oil (stream /$\ and >&> )
is pressurized to 220 kg/cm2 (3100), mixed with hydrogen gas,
and treated in a fixed bed, catalytic reactor at 400 degrees C
(750 degrees F) and 210 kg/cm2 (3000 psig, ref. 5).
Excess flash gas from the hydrotreater is depressurized.
It contains hydrogen, hydrogen sulfide, methane and higher
hydrocarbons. This stream is passed through acid-gas removal
(gas purification* below) and then split. Part of the gas be-
comes product fuel gas ( (£l) and (f£) ); the remainder passes
through hydrogen manufacture (steam reforming) and is re-
cycled to the hydrotreater (ref. 5).
An aqueous bleed stream is also withdrawn from the hydro-
treatment section. This liquor stream, containing ammonia
and H2S, is combined with excess liquor from the oil recovery
section and transferred to wastewater treatment.
(5) Gas Purification
The raw gas from the oil recovery section (stream (jf)
and (7) ) is compressed from 1. 1 kg/cm2 (16 psia) to 28. 8
kg/cm (410 psia) arid mixed with the excess flash gas stream
from the hydrotreatment section. The acid-gas treatment
selected by FMC for these gases is a Sulfinol system consisting
of a solution of sulfolane and diisopropanolamine (DIPA) in
water. This treatment reduces the acid-gas content to 10 ppm
CC>2 and 1 ppm H2S. * The purified gas (stream (fl) and (fy )
can be sold as a low Btu fuel gas, or it can be used to provide
the hydrogen and power requirements of the COED process.
The regenerated acid-gas (stream (g) and (f^> ) goes to the
Stretford sulfur recovery unit.
As discussed in Chapter III, various other acid-gas treatment
processes could have been employed in this application.
VIII-7
-------
(6) Hydrogen Generation
Hydrogen required for the hydrotreatment section can be
produced by either of two methods:
Steam reforming of a portion of product gas
Partial oxidation of product char.
In the former case (illustrated in Figures VIII-1 and VIII-2), the
product gases (stream (fl) and (£|) ) are passed' over a
zinc oxide bed to remove traces of sulfur and thus avoid poison-
ing the reforming catalyst. Next, the hydrocarbon gases are
reformed with steam at 20 kg/cm^ (300 psia) to form carbon
monoxide and hydrogen (CHX + H^O—^ CO + E^). The carbon
monoxide is then shifted with steam to yield carbon dioxide
and additional hydrogen (CO + H2O->CO2 + H2>. Finally,
carbon dioxide is removed by acid gas treatment similar to
that used in the production of fuel gas. The unconverted carbon
monoxide is methanated to prevent deactivating the hydrotreater
catalyst. The resultant hydrogen stream contains approxi-
mately 95- percent hydrogen at 15 kg/ cm^ (200 psia') and
5 percent methane.
In a second method, a portion of the char is reacted with
high pressure steam and oxygen to generate synthesis gas.
The latter is then subjected to quenching to remove fine
particles, a shift reaction to convert carbon monoxide and
water to carbon dioxide and hydrogen, acid gas removal, trace
sulfur removal by a zinc oxide bed, and finally, methanation.
The resultant gas contains again about 95 percent hydrogen.
(7) Steam and Power Generation
It will be assumed here that steam, electric power, and
oxygen requirements of the process are met by on-site plants
fired by a portion of the product fuel gas as illustrated in
Figure VIII-1. A more economical alternative, perhaps less
environmentally acceptable, would be the combustion of the
product char. Combustion of the char from the Utah A Seam
coal yields a flue gas containing 2. 0 kg SO2/106 kcal (1. 1 Ib
SO2/106 Btu), which is below the current maximum permissible
level. The product char from the Illinois No. 6 coal must be
desulfurized before combustion (discussed later).
VIII-8
-------
The oxygen required for the fourth pyrolysis stage is pro-
vided in this analysis by an on-site steam powered oxygen
plant, which produces no harmful emissions.
(8) Energy Balance
The energy balance calculations for the process are pre-
sented in Table VIII-2. In both cases, it is assumed that
process hydrogen is generated by gas reforming. The alter-
native of using char oxidation is calculated for the Utah feed.
The energy efficiencies (coal-to-oil) for conversion of Utah
and Illinois coals to syncrude are approximately 32 percent
and 23 percent, respectively. The overall coal-to-products
thermal efficiencies are about 84 and 79 percent, respectively,
with the remaining energy lost as ambient heat.
(9) Sulfur Balance
The sulfur balances for the two feedstocks analyzed are
presented in Table VIII-3. Note that a significant portion of
the sulfur in the feed coal is retained in the char. For the Utah
coal, 69. 4 percent and 41. 4 percent of the total sulfur is re-
covered in the char when employing gas reforming and char
oxidation respectively. For the Illinois feed, 14.1 percent (ref. 5) of
the sulfur remains in the char. Desulfurization of char from
the high sulfur- Illinois No. 6 coal (discussed later) produces
additional hydrogen sulfide and yields a greater amount of sul-
fur by-products. As much as 84. 2 percent of the sulfur is re-
covered for sale as compared to 29. 4 percent for the Utah coal.
The syncrude produced from Utah and Illinois coals contains
only 0. 02 and 0. 22 percent sulfur (effectively 99 percent sulfur
removal), respectively, and is comparable to natural low-
sulfur, high-grade crude oil. For gas reforming 0. 2 percent
and 0. 8 percent of the feed sulfur is lost to the atmosphere using
the Utah and Illinois feeds, respectively. For the char oxi-
dation mode using Utah coal, 17 percent of the original sulfur
is vented to the atmosphere.
VIII-9
Ai
-------
Table VIII-2
Energy Balance Calculations
Carrier
Coal
Total oner
Syncrude
Gas
Char
Sulfur
Ammonia
Heat Loss
Total encr
Utah A Seam
Heating Value
7,517 kcal/kg
gy input
10.584 kcal/kg
5,374kcal/ in3
7,212 kcal/kg
2.2 13 kcal/kg
5,372 kcal/kg
gy output
Gas Reforming
Amount
378,000 kg/hr
85,405 kg/hr
23,200 m3/hr
187,000 kg/hr
555 kg/hr
869 kg/hr
Energy
(I06kcal/hr)
2,841.4
2.841.4
903.9
1 24.9
1 ,348.6
1.23
4.67
458.1
2.841.4.
(l06Btu/hr)
1 1,275
1 1,275
3,586.8
495.8
5,353.5
4.9
18.6
1.815.4
11,275
Char Partial Oxidation
Amount
378,000 kg/hr
85,405 kg/hr
1 17,100 m3/hr
98,216 kg/hr
771 kg/hr
869 kg/hr
Energy
(I06kcal/hr)
2.841.4
2,841,4
903.9
629.3
708.3
1.71
4.67
593.5
2,841.4
(!06Btu/hr)
11.275
11,275
3,586.8
2,496.5
2,810.5
6.8
18.6
2.355.8
11.275
Carrier
Coal
Total enci
Syncrude
Gas
Char
Sulfur
Ammonia
Heat Loss
Total oner
Illinois No. 6
Heating Value
6.956.1 kcal/kg
;y jnput
10,167 kcal/kg
6.273 kcal/ m3
6.461 kcal/kg
2,213 kcal/kg
5,372 kcal/kg
gy output
Gas Reforming
Amount
378.000 kg/hr
59,61 1 kg/hr
15.930 m-'/hr
208.837 kg/hr
1 1 .760 kg/hr
869 kg/hr
Energy
(I06kcal/hr)
2.629.4
2,629,4
606. 1
99.9
1 ,349.3
26.02
4.67
543.4
2.629.4
<106Btu/hr)
10.433.3
1 0,433.3
2.404.4
395.8
5.354.5
103.3
18.6
2.156.7
10,433.3
VIII-10
-------
Table VIII-3
Sulfur Balance
Carrier
Coal •
Total input
Syncrude product
Char recovered
Sulfur (elemental)
Sult'ides to atmosphere
Sulfur dioxide to atmosphere
Total output
Utah A Seam
Gas Reforming
I03kg/hr
1.89
1.89
0.019
1.312
0.556
0.004
1 .89
tons/hr
2.08
2.08
0.021
1 .446
0.6.13
0.004
2*08
Char Partial Oxidation*
103kg/hr
1.89
1.89
0.019
0.783
0.767
0.008
0.314
1.89
tons/hr
2,08
2.08
0.021
0.863
0.846
0.008
0.346
2.08
Illinois No. 6
Gas Reforming
I03kg/hr
13.98
13.98
0.132
1 .976
11.757
0.1 17
I3i98
tons/hr
15.42
15.42
0.146
2.179
1 2.963
0.129
15.42
This process variation assumes by-product sale of desulfurized fuel gas and use of direct char combustion for
process energy requirements.
3.
DISCUSSION OF CONTROL PROCESSES
The five major waste streams emitted by the COED process
and the proposed control and treatment methods are summarized in
Table VIII-4.
(1) Control of Coal Dust
Most of the coal dust and fines produced during coal
crushing and drying can be controlled with the aid of cyclone
separators and bag filters. The recovered fines can be used
as boiler fuel along with char. Additional precautions, such
as enclosure of coal preparation operations, will minimize
particulate emission (ref. 5).
(2) Utilization of Char
Approximately 50 to 60 percent of the feed coal in the
COED process is recovered as char. Its heating value and
sulfur content are approximately the same as those of the feed
coal, and its volatile content is about 5 percent. This char can
be used as power plant fuel after grinding, though preheated air
VIII-11
-------
Table VIII-4
Nature and Treatment of Major Waste Streams
Waste
Sources
Treatment
Coal dust
Char
Wastewater
Hyrdrogen sulfide
Sulfur dioxide
Coal storage, crushings, drying,
and handling
Pyrolysis reactors, oil recovery,
and filtration
Coal storage, oil recovery, oil
hydro treatment
Raw pyrolysis gas, plant fuel
from hydrotreatment, char
desulfurization, raw gas from
partial oxidation of char
Steam and power plant
Cyclone separators
bag filters, enclosure
Various
Biological and modi-
fied Chevron
Stretford or Claus
plant with tail gas
cleanup
Not required for low
sulfur char or fuel gas
that is acid gas treated
and the addition of a more volatile fuel should be used for light-
off (similar to anthracite combustion).
Upon combustion, char from Utah coal produces a flue
gas containing 2.0 kg SO2/106 kcal (1. 1 Ib of SO2/106 Btu) and
thus does not require desulfurization. Char from Illinois No. 6
coal contains 3. 1 percent sulfur and requires pretreatment to
minimize costly stack gas cleanup. This can be achieved
through:
Gasification to low-Btu gas
Direct desulfurization.
Various procedures applicable for the gasification of char
are covered in some detail in the Synthane process synopsis.
Additional information is presented in Chapter III and in the
process synopses for low-Btu gas production (ref. 5).
VIII-12
-------
FMC has also investigated a process for the direct desul-
furization of COED char. In this approach, the sulfur content of
the thermally stable, highly porous, COED char can be reduced
from 3. 1 to 0. 9 percent by reaction with H2 at 879 degrees C
(1600 degrees F), in the presence of an H^S acceptor, such as
calcined dolomite. The following reactions take place:
Schar + H2 -V U)
CaO + H S »CaS + H O (2)
£i £
+ H0 (3)
CO + H20 ^ C02 + H2 (4)
C + 2H2 ^CH4 (5)
For the production rate of Figures VIII-1 and VIII-2, the de-
sulfurization reactor is a 14. 0 m (46 feet) ID x 17. 0 m (55 feet)
high refractory-lined shaft kiln.
Desulfurized char and spent acceptor are separated and
cooled in a two-stage fluidized bed cooler. At an approximate
fluidizing velocity of 0. 3 to 0. 6 m/sec (1 to 2 ft/sec), the 3 mm
(1/8 inch) acceptor pellets tend to sink to the bottom of the
fluidized bed while the minus 16 mesh char is fluidized above
the acceptor. Acceptor and char are further separated with the
aid of screens and electrostatic precipitators, and the product
char contains about 0. 4 percent dolomite.
Spent acceptor is ground to minus 48 mesh, mixed with
water to form a 39 percent slurry, and regenerated in two 4. 5 m
(15 feet) ID x 30 m (100 feet) high regeneration towers in accor-
dance with the following reactions:
CaS + HO + CO >-CaCO + H S (6)
CaO + CO +• CaCO (7)
£ O
The CO for reactions (6) and (7) could be obtained from the
Stretford off-gas. Reactions (6) and (7) are highly exothermic,
and a temperature of 113 degrees C (235 degrees F) is main-
tained in the regenerator by the evaporation of a portion of
VIII-13
-------
slurry water. Off-gases from the regenerator are sent to the
sulfur recovery plant. The regenerated slurry of calcium car-
bonate and water is dried, the resultant cake is extruded to
form 3 mm (1/8 inch) pellets and calcined in a two-stage cal-
ciner to recover CaO and a COo by-product. The calcined ex-
trusions, heated to 1000 degrees C (1850 degrees F), are re-
cycled to the desulfurizer. Five percent of the acceptor is
assumed to be lost either with char or during scrubbing, drying,
and calcining. The heat required for calcination is provided by
combustion of about 6. 8 x 103 kg/hr (7. 5 short tons/hr) of
product char.
Gasification of char is covered in some detail under the
Synthane process.
(3) Waste Water Treatment (Ref. 5)
The principal liquid waste streams are the leachings
from coal storage, the condensate stream from oil recovery,
and hydrotreatment liquor. The major pollutants in these
streams are coal fines, oil, ammonia, and acid gas. Oils are
removed primarily by an oil water separator, but liquid/liquid
extraction may also be used as an alternative. Although most
oils are recovered, several hundred ppm of hydrocarbons are
retained in the treated wastewater.
The oil-stripped wastewater goes through a Chevron-
type process (in this analysis) for removal of acid gas and
ammonia in two stripping towers. Acid gas is sent to sulfur
recovery; ammonia is recovered as a by-product. The purified
wastewater will still contain the several hundred ppm of oils
and small amounts of ammonia and sulfur compounds. Its
quality is satisfactory for cooling water makeup, but not for
boiler feed water. The low concentration of contaminants in
cooling water are eliminated by biological action in the cooling
system. The blowdowns from cooling towers and boiler can be
treated in a sewage treatment plant, which yields an inert ~
solid residue.
VIII-14
-------
(4) Sulfur Recovery
A Stretford process.was applied in this analysis for re-
covery of elemental sulfur from the acid gas stream of the Utah
A Seam coal because of its low H2S concentration (1. 9 percent).
The process recovers 99 percent of the feed sulfur as an ele-
mental sulfur by-product.
For Illinois No. 6 coal, the H2S concentration of the gas
stream going to the sulfur recovery plant is about 20 percent.
Here, the use of a Glaus plant, followed by the Beavon process
(or other tail gas treatment), is more appropriate. The re-
covery is assumed again to be 99 percent.
(5) Sulfur Dioxide
Stack gas cleanup would not be required for combustion of
char from Utah A Seam coal, because the emissions from
direct combustion would be about 2. 0 kg SO2/10 kcal,
(1. 1 Ib SO?/10 Btu), below the present New Source Performance
Standards.
If char from Illinois No. 6 coal were to be combusted
directly in a steam and power plant, the flue gas would require
stack gas cleanup, before venting to the atmosphere. Alterna-
tively, various lime treatment processes could be employed.
These are considered proven by the EPA.
The Illinois No. 6 coal, after COED lime desulfurization,
will contain about 0. 9 percent sulfur; direct combustion of this
char would emit about 2. 7 kg SO2/106 kcal (1.5 Ib SO2/106 Btu);
therefore, stack gas treatment would still be required,
but at significantly reduced loadings.
4. . COSTS OF POLLUTION CONTROL*'
Costs required for the control and treatment of the pollutants
evolved in the pyrolysis of coal by the COED process are calculated
* The costs presented here are preliminary estimates; the actual
costs of emission control in this process will not be known
until a full-scale facility has been constructed and operated.
VIII-15
-------
in this section. This process is currently at pilot plant development
stage so the cost figures derived are largely based on the scaling up
of the data accumulated during the PDU runs and estimates of costs
for similar pollution control unit processes. The major assumptions
and conventions adopted here are discussed in Chapter III. Alt by-
product flow rate data presented in Figure VIII-1 and Tables VIII-1,
VIII-2, and VIII-3 also apply.
The incremental* capital investment required to control the
pollutants and waste streams discussed in Sections 2 and 3 of this
process analysis are presented in Table VIII-5. The incremental
annual operating costs for these emission control systems are pre-
sented in Table VIII-6. The investment and operating costs required
have been calculated for several feedstock and process alternatives.
The alternative feedstocks analyzed include a low and a high sulfur
coal (0. 5 percent sulfur Utah A Seam coal and a 3. 7 percent sulfur
Illinois No. 6 coal). These reported costs reflect the case where
process hydrogen is produced by reforming a portion of the product
gas into CO + H?. An alternative method of on-site hydrogen generation-
partial oxidation of product char and shifting of the resulting gas is
costed for the Utah coal feed case.
The derivation of the formalae for calculating the incremental
cost of clean fuel production due to pollution control is presented in
Chapter III. For the discounted cash flow (DCF) method of cost
accounting, the required incremental annual cost of clean fuel, X,
for the assumed rate of return is:
X = N + 0.23816 I+ 0.1275 S + 0.230777 W
where
N = incremental net operating cost
I = incremental plant investment
S = startup costs
W = incremental working capital
Incremental costs in this report refer to that portion of the total
plant cost that is assignable to pollution control facilities.
VIII-16
-------
Table VIII-5
(a)
Incremental Capital Investment ($ Million)
Method of Hydrogen Generation
Incremental plant investment '"'
Dust emission control
Wastewater treatment
Stretford sulfur recovery
Claus sulfur recovery with Beavon
tail gas treatment
Char desulfurization
Subtotal plant investment
Project contingencies (c'
Incremental plant investment (I)
Startup costs (d) (S)
Interest during construction 'e'
Incremental working capital ' ' (W)
Incremental capital investment (C)
0.5% Sulfur Utah A Seam Coal
Gas Reforming
1.7
3.7
1.1
—
—
6.5
1.0
7.5 7.5
0.4 0.4
Utility DCF
1.3 1.8
0.2 0.2
9.4 9.9
Char Oxidation
1.7
3.7
1.3
—
—
6.7
1.0
7.7 7.7
0.4 0.4
Utility DCF
1.3 1.8
0.2 0.2
9.6 10.1
3.7% Sulfur Illinois No. 6 Coal
Gas Reforming
1.7
6.5
—
4.5
15.0
27.7
4.2
31.9 31.9
1.1 1.1
Utility DCF
5.4 7.6
0.7 0.8
39.1 41.4
Notes:
(a) Incremental plant investment, return on investment during construction,
and working capital are treated as capital costs in year 0 (the year ending
with completion of startup operation).
(b) Installed costs, including engineering design costs, contractors' profit
and overhead.
(c) Includes costs for unexpected site preparation and hardware requirements
at 15% of plant investment
(e) For the DCF method, computed as the discount rate x incremental
plant investment for 1.875 years' average construction period.
I (l+i)n = I(1.12)'-87S = 1.236761 or 0.236761 additional investment
For the utility financing method, computed as the interest rate
on debt x incremental plant investment x 1.875 years.
(0 Sum of materials and supplies at 0.9% of incremental plant investment
and net receivable at 1/24 of annual incremental revenue.
(d) At 20% of incremental gross operating cost.
-------
Table VIII-6
Incremental Annual Operating Cost ($)
i
oo
Feedstock
Method of Hydrogen Generation
Labor
Direct operating labor (3 men/shift
x $5/hr x 8304 shift hrs/man yr)
Maintenance labor ( 1 .5% of I)
Supervisory ( \S% of direct operating
and maintenance labor
Incremental labor costs
Administration and general overhead
(60% of incremental labor)
Other direct costs*
Supplies
Operating (30% of direct operating
labor
Maintenance ( 1 .5% of 1 )
Incremental cost of supplies
Local taxes and insurance (2.7% of I)
Incremental gross operating cost.
By-product credits
Sulfur (expressed in long tons/day
x 365 days/yr x 0.9 capacity
factor x$IO/LT)
Ammonia (23.0 short tons/day
x 365 days/yr x 0.9 capacity
factor x $25/short ton)
Less by-product credits
Incremental net operating
cost (N)
0.5% Sulfur Utah Coal
Gas Reforming
1 24.600
112.500
35,600
272.700
163,600
1 .000.000
37.400
112,500
149.900
202,500
1.788.700
43.100
188,900
-232,000
1,556,700
Char Oxidation
124.600
115.500
36.000
276,100
165.700
1 .000.000
37.400
115.500
152.900
207,900
1.802.600
59.800
188.900
-248.700
1,553,900
3.7% Sulfur Illinois Coal
Gas Reforming
373,800(9 men/shift)
478.500
1 27,800
980,100
588.100
2,500.000
112,100
478,500
590,600
861,300
5,520,100
912,100
188.900
-1,101,000
4,419,100
•This item includes catalysts and chemicals expended and utilities purchased for the pollution control processes.
-------
For the utility financing case, it is:
X = N + 0.1198 C + 0.0198W
where
C = required incremental capital investment
N and W = as defined above.
The annual production rate of clean fuel products (char, oil and
gas) from the COED process* G, is:
65. 2527 x 1012 Btu/yr (16. 444 x 1012 kcal/yr)
for the Illinois feed and
74. 5795 x 1012 Btu/yr (18. 794 x 1012 kcal/yr)
for the Utah coal. For the char oxidation method of hydrogen gener-
ation using Utah coal, a production rate of
70. 3190 x 1012 Btu/yr (17. 720 x 1012 kcal/yr)
was established.
The incremental cost of clean fuel, due to pollution controls is
found as:
annual cost of product fuel _ _X_
annual production rate G
Solutions to the above equations for the incremental cost of
clean fuel are shown in Table VIII-7. For the various alternatives
costed, the added expense of producing clean energy from the COED
process is found to range from 3. 6 to 14. O^/IO^ Btu of energy pro-
duced.
VIII-19
-------
Table VIII-7
Incremental Cost of Clean Fuel Due to
Pollution Control for the COED Process
Method of Hydrogen Generation
Accounting Method
DCF
Utility inancing
0.5% Sulfur Utah Coal
Gas Reforming
$/yr
3,440,100
2,686,800
$/106Btu
0.046
0.036
Char Oxidation
$/yr
3,484,900
2,708,000
$/106 Btu
0.050
0.039
Method of Hydrogen Generation
Accounting Method
DCF
Utility financing
3.7% Sulfur Illinois Coal
Gas Reforming
$/yr
12,341,500
9,117,200
$/106Btu
0.189
0.140
VIII-20
-------
5. REFERENCES
(1) Jones, J.F., "Project COED," Paper presented at
Institute of Gas Technology Symposium on Clean Fuels
from Coal, Chicago, 10-14 September 1973.
(2) "Char Oil Energy Development, " Project COED (FMC),
Final Report, Office of Coal Research, Department of
the Interior, June 1971.
(3) "Char Oil Energy Development, The Desulfurization of
COED Char, Part III, " Project COED (FMC), Interim
Report No. 2, Office of Coal Research, Department of
the Interior, May 1970.
(4) "Char Oil Energy Development, " Project COED (FMC),
Interim Report No. 1, Office of Coal Research, Depart-
ment of the Interior, February 1970.
(5) Specific Comments from Process Developer, Correspon-
dence from FMC Corporation, April 1974.
1. Product Yield
Yields shown in Table VIII-1 were obtained from
published data. COED pyrolysis yields are now as
follows:
Utah Illinois
A Seam No. 6
Char, ton/ton dry coal 0. 55 0. 54
Oil, bbl/ton dry coal 1.32 1.09
Pyrolysis gas, MSCF/ton 10.3 10.4
2. Process Description
Combustion of char must be with steam-oxygen
mixture to supply heat requirements for the
last three stages. Using oxygen only would
cause localized overheating and clinkering.
The hydrogen produced using steam-oxygen
VIII-21
-------
for the heat requirements for the last three
stages is used to produce all the hydrogen re-
quired in the hydrotreating section of the plant.
This eliminates the need of char oxidation as a
source of hydrogen generation.
be done at pressures rang-
4-^ OAAH +~n-tAr+
Hydrogenation may be done at p]
ing from 1500 psig to 3000 psig.
High pressure separation is used to separate
the oil and hydrogen rich gas stream. The
hydrogen rich gas stream is recycled to the
reactor through a recycle compressor. A
small purge stream of the recycle gas stream
is sent to the purification section for use in
hydrogen manufacture.
After pyrolysis, the sulfur remaining in the
char is about 45 percent of the sulfur in the
Illinois feed, not 14.1 percent as indicated.
3. Discussion of Control Processes
Coal dust and fines generated in the process
are fed into the fourth stage reactor and gasi-
fied. The filter cake from the filtration sec-
tion is also fed into the fourth stage reactor.
All wastewater streams from the COED pro-
cess are recycled to the fourth stage reactor.
This destroys all the organic contaminants
and greatly reduces the wastewater treatment
required for a COED facility.
4. Additional Comments
The use of char desulfurization has been recom-
mended to meet environmental restrictions with
high sulfur coals. However, the char cannot be
sufficiently reduced in sulfur level, and stack gas
treatment is still required to remove the SO2 after
char combustion. The use of both char desulfuri-
zation and stack gas scrubbers may not be the best
VIII-22
-------
choice. Char gasification is a preferred way of pro-
ducing a low sulfur, environmentally acceptable fuel
gas. This approach is noted in the report but not
discussed in detail.
VIII-23
-------
IX. THE SOLVENT REFINED COAL (SRC) PROCESS
The Solvent Refined Coal (SRC) process is essentially a mild
hydrogenation process that produces a low ash, low sulfur, high heating
value liquid or solid fuel from coal. It originated as a result of re-
search performed from 1962 to 1964 by the Spencer Chemical Company
on a grant funded by the Office of Coal Research (OCR). In October
1966, OCR contracted the Pittsburg and Midway Coal Mining Company
(PAMCO), an affiliate of Spencer Chemical, to design, construct and
operate an SRC pilot plant. This plant used various grades of bitumi-
nous and lignite coals.
In the SRC process, a highly aromatic solvent, a by-product of
the process, dissolves coal at 454 degrees C (850 degrees F) and
71. 3 kg/cm^ (1000 psi). The solution is filtered to remove coal ash
and unreacted carbon, and the solvent is removed by distillation. The
coal product contains substantially less sulfur than originally con-
tained in the feed coal and little or no ash. It is, therefore, an
attractive fuel for power plants and other industrial uses. The de-
ashed coal product also appears to be an attractive feed for delayed
coking units which make electrode grade coke*.
The solid coal product may be further hydrotreated to make
liquid hydrocarbons. This process is similar to the hydrocracking of
petroleum feedstocks and is discussed in the Gas Combustion Retort
and COED process synopses.
PAMCO is constructing a 45 m ton/day (50 short tons/day) pilot
plant, designed by Stearns-Roger Corporation at a site in Taconia,
Washington. This $17. 5 million plant is now nearing completion.
1. PROCESS DISPLAY
The flow diagram and energy balance for the SRC process is
presented in Figure IX-1, and Tables IX-1, IX-2 and IX-3 show the
Electrodes made of high purity coke are used in many chemical
processes.
IX-1
-------
FIGURE IX- 1
The Solvent Refined
Coal Process
><
(O
^T~ CLAUSA
SCOT SUL
RECOVE
t_ 1
*— lo.ioi—
00
so,
3.77!
> r
^""1 1
UR I
'" 1 1
1
J
[9.871
ENERGY BALANCE
CARRIER
COAL
NATURAL GAS
TOTAL INPUT
SYN-COAL
SULFUR
PHENOLS
CRESYLIC ACID
LIGHT OILS
ELECTRICITY
COOLING WATER
OTHER
TOTAL OUTPUT
X 106 ktn,h,
2.692.4
194.3
2,886,1
1.940.6
21^
ia?
35.6
212.4
J7J
674.2
13.6
2.888.7
132.4 MW t>porud>
•StMcific mnrrwnt wid
-------
Table IX-1
Composition of Gaseous Streams
Streams
Components, mole %
H2
CO2
H2S
S02
N2
Cj
C2
C3
C4
H2O
02
Total
(Ft3/hr)x 106
(Cubic meters/hr) x 104
Lb-Moles/hr
(Gm-moles/sec)
Light oil Ib/hr
Wash solvent, Ib/hr
(kg/hr)
Process solvent, Ib/hr
(kg/hr)
Phenols, Ib/hr
(kg/hr)
© ©
100.00 36.73
— 2.30
4.38
— —
— —
43.88
9.50
.2.85
0.36
—
100.00 100.00
2.06 1.33
5.83 3.77
5,448 3,500
(687) (441)
1 101
499
©
22.24
1.49
3.15
27.54
6.63
2.40
0.42
36.13
—
100.00
5.34
15.12
14,094
(1,777)
25 138
11 403
108,558
49,242
3,000
1,361
©
36.73
2.30
4.38
—
—
43.88
9.50
2.85
0.36
—
—
100.00
3.22
9.12
8,499
(1,072)
2 617
1 187
©
36.73
2.30
4.38
—
—
43.88
9.50
2.85
0.36
—
—
100.00
1.89
5.35
4,999
(630)
1 516
678
®
14.85
2.83
11.71
40.84
19.15
9.12
1.50
—
—
100.00
0.62
1.76
1,644
(207)
1 261
572
©
1.23
1.14
19.17
—
9.12
23.07
35.22
11.05
—
100.00
0.03
0.08
82
(10)
678
308
0
30.95
2.42
6.35
42.71
12.02
4.78
0.77
—
—
100.00
2.55
7.22
6,724
(848)
3 455
1 567
© ® © ©'
33.92
lOppmV 27.55 15.99
SppmV 72.45
2.27
79.00
46.81
13.18
5.24 - -
0.85
100.00
2.74
100.00 100.00 100.00 100.00
2.33 0.22 4.34 1.14
6.60 0.62 12.29 3.23
6,135 589 11.443 3,013
(774) (74) (1,443) (380)
3/155
1 567
© ©
9.10
2.14
16.98
79.00
33.01
21.02
13.71
4.04
21.00
100.00 100.00
0.12 4.34
0.34 12.29
308 11,443
(39) (1,443)
19 890
9 022
47/117
21,327
QS 914
©
16.55
3.55
2.16
75.14
2.60
100.00
4.56
12.91
12,032
(1,517)
•Only the water vapor content of this stream can be estimated from the source data.
-------
Table K-2
Composition of Liquid Streams
Stream
Components, mole %
H-,
CO,
c,
c,
c3
C4
Total
Lb-Moles/hr
(Gm-moles/sec)
Light oil, Ib/hr
(kg/hr)
Phenols, Ib/hr
(kg/lir)
Cresylic acid. Ib/hr or kg/hr
Wash solvent Ib/hr
(kg/hr)
Process solvent. Ib/hr
(kg/hr)
(kg/lu)
(kg/hr)
Water Ib/hr
(kg/hr)
Feed coal Ib/hr
(kg/hr)
A A A
18.20
2.60
11.02
41.34
16.44
8.37
2.03
100.00
1.275
(161)
31,295
14,195
1,666,667 1,666,667 1,558,109
756,000 756,000 706,758
488,376
TM.527
90025
48,835
81* 1"
378,000
A
8.80
2.11
17.24
32.44
21.09
14.04
4.28
100.00
318
(40)
24,462
11,096
106,979
48,526
1,535,074
696,310
488 376
221,527
A
1.04
24.93
15.90
23.23
23.70
11.20
100.00
11
(1)
4,572
2,074
59,962
27,199
1,325,166
601,095
A
0.81
0.67
14.02
4.98
16.33
40.23
22.96
100.00
123
(15)
47,983
21,765
54,797
24,856
341,501
154,905
•
A A A
0.75
0.70
14.89
5.85
16.88
38.91
22.02
100.00
134
(17)
52,555
23,839
114,759 114,759 94,025
52,055 52,055 42,650
1,666,667
756,000
A A A
2.59
2.87
14.33
30.26
25.42
18.93
5.60
100.00
503
(63)
22,521
10,216
3,000
1,361
208,784
94,704
108,558
49,242
91.667
A
8.17
0.73
7.16
44.70
39.24
100.00
52
(6)
51,877
23,531
/14\ /1S\
21.33
2.76
8.94
44.30
14.89
6.49
1.29
100.00
956
(121)
6,833
3,099
9
4
lOOppmW
7,780
3,529
23,035
10,449
9l,fifi7
A
—
—
—
—
—
—
—
8,427
3,822
X
-------
Table IX-3
Composition of Solids Streams
Stream
Components, weight %
C
H
N
S
O
Ash
Mositure
Total
Coal, Ib/hr
(kg/hr)
Precoat, Ib/hr
(kg/hr)
Mineral residue Ib/hr
(kg/hr)
Ash, Ib/hr
(kg/hr)
Heating value, Btu/lb
(kcal/kg)
DO
65.80
4.36
1.00
3.14
9.56
6.63
9.51
100.00
896,000
406,426
--
--
--
11,924
6,624
H]
--
--
250
113
--
--
--
--
--
--
GO
27.01
--
9.60
--
63.69
--
100.00
--
--
250
113
90,025
40,835
--
--
4,191
2,328
S
3.95
--
--•
0.54
95.51
--
100.00
--
--
250
113
--
59,750
27,103
--
--
[E
88.16
5.23
1.54
1.17
3.42
0.48
--
100.00
488,376
221,527
--
--
--
--
15,768
8,760
IX-5
-------
gaseous, liquid, solid stream compositions and flow rates referred to
in Figure IX-1. The data used in this analysis are based on an economic
evaluation performed by Stearns-Roger Corporation for OCR (referenced
at the end of this chapter) completed in 1969. At this writing more
recent data were not available; however, it is very likely that the
process has been further developed in the past four or five years.
(1) Bases for Analysis
The analysis shown is for a 3. 14 percent sulfur, Kentucky
No. 11 coal whose composition is shown in Table IX-4.
Table IX-4
Component Analysis for Coal Feedstock*
Component
Carbon
Hydrogen
Nitrogen
Sulfur
Oxygen
Ash
Moisture
TOTAL
Weight %
65.80
4.36
1.00
3.14
9.56
6.63
9.51
100.00
. "Kentucky No. 11
This plant converts 9,754 m tons/day (10,752 short tons/
day) of coal feed to 5, 316 m tons/day (5, 860 short tons/day) of
a depolymerized, de-ashed (0. 48 percent ash) coal product
containing 1.17 percent sulfur. The heating value of the solvent
refined coal of 8, 760 kcal/kg (15, 768 Btu/lb) equates to a daily
plant output of 46, 620 x 106 kcal (185 x 1Q9 Btu). The plant's
thermal efficiency, coal plus fuel gas feed-to-coal product, is
determined to be 67 percent.
(2) Layout and Symbols
The general direction of the process flow is from the coal
supply on the left to the product fuel on the extreme right of the
flowsheet, with residuals and by-products shown along the
IX-6
-------
bottom. The bold line indicates the flow of the primary dissolu-
tion process. The portion of the process contained within the
dashed lines was not defined by the process developer but was
assumed for this analysis.
Streams labelled with circles, triangles or rectangles refer
to the gaseous, liquid or solid stream compositions shown in
Tables IX-1, IX-2 and IX-3, respectively. The overall energy
balance is shown on Figure IX-1 and calculated in Table 1X-5.
The rhombic-shaped units represent intermediate products
or uses for which the distribution is not shown. Pollutant cleanup
processes are indicated by sloping rectangles. The treatment
and recovery processes which this symbol represents are dis-
cussed in the process description and pollution control sections
of this process summary. For an in-depth explanation of each
licensed process mentioned, refer to Chapter III as well as to
readily available process literature and to the bibliography in-
cluded in this report.
2. PROCESS DESCRIPTION
The unit processes which define the SRC process are described
in this section. By-product manufacture and pollutant generation are
discussed where appropriate. The discussion of pollution control
processes is covered in the next section.
In brief, the feed coal is crushed, dried and mixed with a
process-derived solvent. The mixture is preheated and dissolved in
a hydrogen atmosphere. Ash and unreacted carbon are filtered from
the resulting liquid before it is flashed (rapidly vaporized) to remove
solvent and light oil fractions containing sulfur. The remaining liquid
is solidified by cooling, yielding the low sulfur product fuel.
It should be stressed that the characteristics of the solid, coal
product from the SRC process is a function of the feedstock selected.
In particular, the sulfur content of the product will be a function of
the quantity and character of the sulfur in the raw feed coal. An
improvement in desulfurization is expected if the sulfur content of
the feed is primarily pyritic rather than organic since the decomposed
pyritic-sulfur (ferrous sulfide) can be readily removed with the ash.
Similarly, coal with high ash alkalinity will tend to remove the sulfur
as calcium sulfide in the ash. The degree of desulfurization of the
pyritic and organic sulfur to H^S gas is a complex function of time,
temperature and pressure which must be determined experimentally.
IX-.7
-------
Table IX-5
Energy Balance Calculations
Carrier
Calculations
Energy
(xl06Btu/hr)
x!06kcal/hr
Coal
Natural Gas
Total energy input
896,000 Ib/hr x 11,924 Btu/lb
10,684.2
771.0
11,455.2
2692.4
194.3
2886.7
i
oo
Solvent refined coal
Sulfur
Phenols
Cresylic acid
Light oil
Electricity (exported)
Cooling water
Other* •
Total energy output
488,376 Ib/hr x 15,768 Btu/lb
10.88 short tons/hr x 2000 Ib/short tons x 3783.4 Btu/lb
1.5 short tons/hr x 2000 Ib/short tons x 14,117.6 Btu/lb
5.0 short tons/hr x 2000 Ib/short tons x 14,117.6 Btu/lb
22.18 short tons/hr x 2000 Ib/short tons x 19000 Btu/lb
32,390 kW x 3413 Btu/kWh
165,000 gal/min x 60 min/hr x 8.34 Ib/gal x 30°F A T x
1 Btu/lb°F
By difference
7,700.7
86.7
42.4
141.2
842.8
110.5
2,477.0
53.9
11,455.2
1940.6
21.8
10.7
35.6
212.4
27.8
624.2
13.6
2886.7
*Include heating values of ot.ier products, sensible heat of product streams, heat lost to the atmosphere.
-------
If the sulfur content of the depolymerized coal product is too
high for direct combustion, as in the case presented in this synopsis,
the solid product may be further hydrodesulfurized to yield a fuel oil
product. Alternatively, stack gas from the combustion of this coal
can be desulfurized more readily than if raw untreated coal had been
burned.
(1) Coal Preparation
The coal is mined, crushed to 2. 54 cm (1 in) and trans-
ported to the plant by conveyor belt, truck or train. Suitable
coal storage must be provided at the mine and at the plant. At
the plant, the coal is further crushed to 0. 3 cm (. 125 in) in
double-ring disintegrator-type crushers by recirculating the
oversize material (20-30 percent of the load). The oversize
product is separated over screens for recirculation to the
crushers. Two parallel circuits of grinding equipment are re-
quired due to the capacity limitations of the disintegrating equip-
ment. Screen undersize drops into a distributing bin. Coal is
conveyed from this distribution bin to flash dryers, also con-
tained within the coal preparation section shown in Figure IX-1.
The coal passes countercurrent to hot flue gas from the direct
fired preheaters just upstream of the dissolvers. The flash
dryer exhaust gas (stream (£2) ) is cleaned in mechanical flue
gas cleaners and precipitated dust is returned to the coal stream.
(2) Slurry Preparation
The feed processed through the coal preparation section
of Figure IX-1 is weighed into slurry mix tanks to control the
solvent-to-coal ratio. This wash solvent is generated con-
tinuously, and recirculated through the process. A centrifugal
recirculating pump is used to mix the slurry. Since the coal
leaving the preparation section is only partially dried, residual
moisture will be vaporized by the hot solvent in the flash tanks
and vented through moisture condensers provided on the slurry
mix tanks (stream fl.6\ ).
IX-9
-------
(3) Preheater Section
Dissolving of the coal takes place in two steps in the SRC
process: the preheater section and the coal dissolution
section. In the preheater section, the coal-solvent slurry is
exposed to hydrogen under controlled time, temperature and
pressure conditions. It is heated to dissolution temperature
before flowing to the coal dissolution section where the primary
solation occurs.
Reciprocating pumps are used to feed the slurry to the pre-
heaters since multi-stage centrifugal pumps, for transporting
the 0. 3 cm (. 125 in) coal particles in hot solvent, are not readily
available. Multiple pumps are provided to proportion the slurry
among the 16 streams that are required in this section. These
streams are joined by a measured amount of hydrogen and then
further divided into 32 passes through four preheaters so as to
maintain a low fluid velocity in the pipes and thus minimizes
erosion, to maintain economically small pipe diameters, and
to assure the required time-temperature profile. A steam coil
in each preheater recovers waste heat from the stack gases
before they are transferred to the coal preparation area where
it is used to dry the feed.
(4) Coal Dissolution Section
Four dissolvers are used for this facility design. At a
pressure of 71. 3 kg/cm^ (1000 psig) and temperature between
440 degrees and 455 degrees C (825 degrees to 850 degrees F),
a retention time in the dissolvers of 15 minutes is adequate to
dissolve most of the organic matter in the coal. Dissolver
effluent is cooled to 330 degrees C (about 625 degrees F) where
its liquid and vapor phases, stream^2\and(3)respectively are
separated. The liquid slurry stream is sent to the filtration
and separation system. The vapor stream proceeds to a com-
plex waste heat recovery system where it is partially condensed.
It is finally cooled to 43 degrees C (110 degrees F) in a cooling-
water heat exchanger. This stream is separated into three
streams by a high pressure condensate separator:
Stream /i 2\ — an aqueous stream which is directed
to the phenol and cresylic acid recovery unit
IX-10
-------
Stream A JV — a hydrocarbon stream proceeds to a
complex flash system where it is depressurized to
about 12 kg/cm2 (155 psig). The flashed vapors,
stream (6) , are sent to the acid-gas treatment
section while the remaining liquid, stream/6\.
proceeds to solvent recovery (discussed later)
Stream® — a high pressure, 43 degrees C (110
degrees F) vapor stream which is further divided
into two fractions:
One stream (stream (D ) is recompressed,
mixed with makeup hydrogen, and fed to the
slurry entering the preheater
The other stream (stream (5) ) is expanded
and sent to the acid-gas treatment section.
Each of these streams is discussed further when the appropriate
processing units are considered.
(5) Filtration and Separation System
The liquid stream from the coal dissolution section pro-
ceeds to the filtration and separation system where ash is
removed. Experts in the field of solid-liquid separation agree
that the problem presented by this SRC process is exceptionally
difficult.
The Consolidation Coal Company evaluated three solid-
liquid separation techniques in its coal liquification pilot plant:
filters, centrifuges and cyclones. They found cyclones to be
the most reliable. It should be noted, however, that the filters
at this plant were purchased on a least-cost basis; they were
the first units of this type made by the manufacturer; and, they
suffered from inherent mechanical problems. Since the product
fuel generated by the SRC process may eventually undergo
catalytic hydrocracking to produce a syncrude, a solid-free
stream is necessary. This requirement eliminates the use of
centrifuges or cyclones, leaving filtration as the primary
technique available. This technique has, therefore, been
selected for discussion in this analysis.
IX-11
-------
The solvent and dissolved coal slurry, stream^is filtered
through rotary precoat filters containing a bed of diatomaceous
earth. The filters are enclosed in a pressure vessel maintained
at 11.6 kg/cm2 (150 psig) and 316 degrees C (600 degrees F).
The solids entrained in the filters are washed by a light solvent
(stream /.0\ ) from the fractionation section (to be discussed)
and are removed to the mineral residue processing zone.
In this type filter, solid-liquid separation is accomplished
by forcing the slurry through the layer of diatomite precoat
previously deposited on the exterior surface of the drum-type
filter. The resulting filtrate is essentially free of solid. About
100 ppm of suspended material may remain. After spraying a
light solvent on the filter to wash out the residual filtrate, the
outside layer of diatomite is shaved from the precoat to ensure
a fresh surface for further filtering. About 50 to 100 hours of
operation are expected before the precoat is exhausted and a
new layer of diatomaceous earth must be applied. As seen from
Table IX-3, stream[|], about 113 kg/hr (250 Ib/hr) of precoat
material will be consumed in this operation. The shavings join
the filtered solids and some wash solvent in the mineral residue
processing zone.
(6) Mineral Residue Processing Zone
In the mineral residue processing zone, the filter
shavings and wash solvent, containing minimal coal-oil, drop
into a pressurized surge pot which is fed to dryers by a screw
conveyor. Screw conveyors are used because of the gummy
characteristics of this stream. The dryers, rotary drum type,
recover the wash solvent (stream £&) which is reused in the
filtration system.
The remaining dried solid, stream [3] , containing carbon,
coal ash and 9. 6 percent sulfur is sent to the steam and power
generation section. Alternatively this residue may be oxygen
gasified to produce fuel gas for steam generation and hydrogen,
eliminating the need for natural gas supply.
(7) Steam and Power Generation
The solid stream (stream \3] ) from the mineral residue
processing zone is used as fuel in fluidized bed boilers(or after
IX-12
-------
gasification in direct-fired boilers) to generate the steam and
electrical energy requirements of the process. The ash (mineral
residue remaining after carbon has been consumed during com-
bustion) can be stored in silos and periodically transferred to the
worked-out mining areas for disposal. The steam generated is
saturated at a pressure of 15 kg/cm2 (200 psig). The sulfur con-
tent of the flue gas is sufficiently high that it must be further pro-
cessed to minimize air pollution.
This process, as designed, makes significant use of waste
heat boilers to recover process steam. Sufficient steam is
generated to supply over one million pounds per hour (about
450,000 kg/hr) for process steam or for generation of electricity.
The electrical generating capacity of the facility is sufficient to
generage 32. 4 MW of electricity more than the 29. 5 MW needed
to satisfy process requirements.
(8) Solvent Recovery System (Vacuum Flash, Mash System,
Fractionation System)
The filtrate from the filtration and separation system
(stream,/^) contains the solubilized coal, the slurry solvent
and some of the wash solvent used in the filtration unit. The
filtrate is first preflashed to remove the lightest materials.
This vapor (stream (T3) ) is combined with the condensate from
the high pressure condensate separator (stream fl.r\ ). This
combined stream is flashed through intermediate and low
pressure flash vessels in the flash system where some of the
H2S is removed (stream (6)). The remaining liquid is pre-
heated and sent to the fractionation system (stream/p\). The
liquid from the preflash unit is next flashed in the vacuum flash
vessel where the overhead vapor (stream /5\) is removed. This
steam is condensed with cooling water and then pumped to the
fractionation system also. The liquid remaining in the vacuum
flash vessel is sent to the product solidification section.
(9) Product Solidification
In this section, the hot, stripped and de-ashed product
from the vacuum flash vessel, a liquid at this elevated tem-
perature (200 degrees to 260 degrees C or 400 degress to 500
degrees F), is cooled in drums which solidifies the product.
IX-13
-------
This cleaned coal can then be readily transported on belt con-
veyors to storage bins for subsequent rail shipment to
purchasers.
With further hydrogenation, such as by low-temperature
catalytic cracking, this solvent refined coal can be converted
into hydrocarbons which are liquid at room temperature.
(10) Hydrogen Generation
The SRC process includes a conventional plant for
generating hydrogen by the steam reforming of natural gas
(CH4 + H2O—>3H2 + CO) followed by shift conversion to yield
a hydrogen-rich gas stream (CO + H2O-> H2 + CO2). This
process of on-site hydrogen generation is discussed in more
detail in Chapter XI on hydrodesulfurization of fuel oil. This
plant must satisfy process requirements for 1.4 x 106m3/day
(50 x 10"ft^/day) of hydrogen. The alternative use of mineral
residue through gasification to produce fuel gas and hydrogen
would eliminate the need for a natural gas reformer.
(11) Energy Balance
The overall energy balance for the SRC process is pre-
sented in Figure IX-1 and derived in Table IX-5. From this
energy balance, the thermal efficiency of the process is found
to be:
7, 700. 7 x 106Btu/hr of cleaned product
coal -f-11, 455. 2 x lo6Btu/hr of feed =
67. 2 percent
(12) Sulfur Balance
The sulfur balance for the 3.14 percent* sulfur, Kentucky
No. 11 coal feedstock is quantified in Table IX-6 and discussed
in Section 3 of this chapter. The reported sulfur output of
12, 773 kg/hr (28,160 Ib/hr) is the total amount of sulfur evolved
(220 Ib/hr) vented to the atmosphere from the sulfur recovery
plant, 64 kg/hr (140 Ib/hr) remaining with the liquid by-products,
and 2594 kg/hr (5720 Ib/hr) remaining in the solvent refined
product coal.
3. 74 percent by weight when calculated on a dry, ash-free basis.
IX-14
-------
Table IX-6
Sulfur Balance for the SRC Process
Carrier
Input
Coal feed
Total input
Output:
Solvent refined product coal
Elemental sulfur
Ash
Discharged to atmosphere
Liquid by-products
Total output
Ib/hr
28,160
28,160
5,720
21,760
320
220
140
28,160
kg/hr
12,773
12,773
2,594
9,870
145
100
64
12,773
3. DISCUSSION OF POLLUTION CONTROL PROCESSES
The major waste streams occurring in the SRC process and
their proposed treatment are discussed in this section and summarized
in Table IX-7.
(1) Acid-Gas Removal
Several of the sulfur containing gas streams from the
process are combined and sent through an MEA (monoethanola-
mine) acid-gas treatment facility to recover the sulfur in the
form of H2S. These streams consist of the
Overhead vapor stream from the high pressure
condensate separator (stream(5)) .
Effluent from the flash system (stream(?))
Overhead gases from the fractionation system
IX-15
-------
Table IX-7
Source and Treatment of Major Waste Streams
Waste
Source
Treatment
Coal dust
Hydrogen sulfide (intermediate
stream) and sulfur dioxide
Phenols, cresylic
Acids, light oil
Mineral residue
Combustion gases
Crushers, driers
MEA acid-gas treatment
and fluidized bed boilers
Water condensate
Fluidized bed boilers
Process heaters
Cyclone separators,
bag filters
Claus and SCOT
Solvent extraction,
active carbon adsorption
None required. Used as minefill
None required. Vented to atmosphere
Though the process developer did not specify the type of acid-
gas removal, it was assumed that a simple MEA system would
be satisfactory. The primary difficulty with MEA systems is
that they do not handle carbonyl sulfide (COS). Due to the high
water vapor-to-CC>2 ratio in this stream, however, COS is not
expected to be present in any significant quantity*. The de-
sulfurized gas (stream (9)) contains propane, butane and some
light oil. It can be readily consumed in direct fired heaters and
preheaters to generate process steam. The flue gas, containing
only 8 ppmV of sulfur requires no further treatment and can be
vented to the atmosphere.
The sulfur-rich regenerator effluent stream (stream @ )
contains about 72 percent H2S. This is satisfactory feed for a
Claus plant to produce elemental sulfur; however, the process
developer has added a further complexity as discussed in the
next subsection. If the mineral residue is gasified, additional
sulfur removal will be required for that stream.
(2) Sulfur Recovery
The process developer suggests combining the concentrated
H2S-rich stream from the acid-gas treatment section (stream @
with stream (fi), the SO2 containing flue gas from the steam and
power generation section. This combined stream is used as the
feed to the sulfur recovery plant.
The quantity of CO2 and H2S present in the reaction, COS -f-
CO2 + H2S implies negligible sulfur existing as COS.
TX-16
-------
The amount of sulfur dioxide generated during steam and
power generation is approximately correct for oxidizing the
quantity of hydrogen sulfide recovered from the acid-gas treat-
ment unit. Conceptually, therefore, these streams could be
mixed and reacted over Claus catalysts to form elemental
sulfur. However, the effective sulfur concentration feeding
this Claus unit is equivalent to only a 6. 5 percent H^S stream
feeding a conventional Claus plant. With this low concentration
stream, the efficiency of the Claus catalyst is reduced and
sulfur recovery, even with multiple stages, is not expected to
be over 50 to 70 percent. A tail gas purification unit, therefore,
also becomes a requirement.
The SCOT process, developed by the Shell Development
Company, could be applied to treat the Claus tail gas. In the
SCOT process, the Claus feed is hydrogenated over cobalt-
molybdenum catalyst which reduces the sulfur compounds to
H2S. This H2S is selectively withdrawn by a diisopropanolamine
(DIPA) system*. The H^S stream from the DIPA system is
recycled to the Claus plant, thereby improving the H2S-to-
SO2 ratio in the Claus feed.
Engineering practice suggests that the modified Claus
sulfur recovery facility, as discussed above, might not be the
most desired alternative. During plant upsets, the flow, and
therefore the ratio of the E^S to the SO2 may vary. In addition,
when processing different coals, this critical H^S : SO2 ratio
may be upset due to the varying proportions of organic and
pyritic sulfur in the feedstock. It is, therefore, suggested that
these two streams be processed separately as follows:
The H2S-rich stream is an excellent feed for a con-
ventional Claus plant. Any Claus tail gas purifica-
tion process can be added.
The flue gas stream from the steam and power
generation section contains about 2 percent SO2.
This is too high for most direct stack gas cleanup
systems. An operable scheme might be a Wellman-
Lord process to concentrate the SO2 followed by an
Allied process to reduce it to elemental sulfur. In
A DIPA system could have been chosen for the main acid-gas
removal unit so that a single regenerator could have been used
for both the acid-gas plant and the Claus tail gas plant.
IX-17
-------
this scheme, one Glaus plant could be used for both
the acid-gas stream and the effluent stream from
the Allied process, but these streams would be in-
dependent so that plant upsets should not significantly
affect sulfur emissions. If the mineral residue is
gasified, a separate SC>2 stream is not generated
and the Glaus plant operation is simplified.
(3) Solid Waste Streams
Coal dust is generated during the crushing and drying
processes of the feed preparation step. Cyclone separators are
used to remove most of the fines. Bag filters complete the
cleanup before the gases are vented to the atmosphere. These
collected solids are directed to the feed slurry tank for use in
the process. The quantity of fines expected to escape the dust
removal system is minor.
The largest volume of solids from the plant is the mixture
of coal ash and filter material which is removed from the
fluidized bed boilers (explained in Section 2). Because this
material has been well burned, it contains a minimum amount
of carbon and a low concentration of sulfur. Residual solids
from the fluidized bed boilers are sent to the mining area for
disposal. Solid residue from the wastewater treatment section
is also sent to the worked out portions of the mines. All of this
material is relatively inert and should pose no problem of
leaching or oxidation.
Periodically, the catalyst in the shift conversion reactor
of the hydrogen plant must be replaced. A conventional shift
conversion catalyst is iron oxide. Before removal from the
reactor vessel, the catalyst is carefully oxidized in air to pre-
vent auto oxidation. After removal from the reactor vessel,
the catalyst is also discarded in the worked out mining area,
(4) Liquid Waste Streams
The liquid waste streams from this facility will contain
phenol and cresylic acids. * More development work is required
An oily mixture of phenol and cresol compounds.
IX-18
-------
to demonstrate a process for the recovery of cresylic acids
from water streams, although the technology for phenol re-
covery is well known. In this analysis, the cresylic acids are
reduced to about 100 ppm in the wastewater by a combination of
solvent-extraction and stripping. Cresylic acid is then recovered
from the solvent by distillation or a combination of distillation and
extraction. Process water, containing residual phenol and cresylic
acids, is filtered through activated carbon. The carbon absorbs
cresylic acids and phenol. These compounds are then regenerated
by stripping with 15. 1 kg/cm^ (200 psig) steam. The stripped
acids are reprocessed in the recovery plant. An alternate to
recovering the phenols and cresylic acids would be consumption
of the wastewater with destruction of these components in the
mineral residue gasifier.
Although not discussed by the process developer, it has
been calculated that 790 kg/hr (20. 9 short tons/day) of anhydrous
ammonia can also be produced from the aqueous stream from the
high pressure condensate separator (stream /.2\ ). To recover
this ammonia, after phenol and cresylic acid separation, the
wastewater must be passed to a Chevron unit for further purifi-
cation. Some H^S and CO2 are steam-stripped and sent to the
acid-gas treatment unit. The remaining anhydrous ammonia
can be recovered for sale.
Process water runoff goes to a separate storm sewer
system and to holding ponds for further treatment. Skimming
and aeration will constitute sufficient treatment for this water.
Runoff from other areas will flow into the storm sewage system
for discharge without treatment.
Cooling water treatment for corrosion reduction will use
nonpoisonous compounds such as organic zinc borates and will
not use chromates. Cooling tower blowdown will flow to the
sewer without further treatment. Sewage will be treated by
standard methods to give an effluent of acceptable quality arid a
biological oxygen demand for discharge to the watershed.
(5) Waste Gases
The SRC process, as designed, employs a modified Glaus
unit to convert the sulfur in the SC^-containing flue gases from
DC-19
-------
the mineral residue fuel boilers, and H2S from the acid-gas
removal section. The quantities of SO2 and H2S are in nearly
stoichiometric ratio (2H2S + SO2 -* 2 H2O + 3 S) and therefore
no conventional reaction chamber is included in the Glaus plant.
Two stages of reactors are used, followed by a SCOT tail gas
cleanup system. The overall sulfur recovery from the plant is
expected to exceed 99 percent of the entering sulfur, with less
than 1 percent of the sulfur entering the plant in the coal feed
being vented to the atmosphere.
4. COSTS OF POLLUTION CONTROL*
The costs of the pollution control and treatment processes de-
scribed in Sections 2 and 3 are calculated here. Since the SRC process
for producing a syn-coal product is not yet in commercial use, the cost
figures used here are estimates of costs for similar pollution control
unit processes and from extrapolating SRC process pilot plant data.
The major assumptions and conventions adopted here are discussed
in Chapter III of this report. All by-product flow rates appearing in
Figure IX-1 also apply.
The incremental** capital investment for the control of the po-
tential pollutants discussed in this synopsis is presented in Table IX-8.
The incremental annual operating costs appear in Table DC-9.
As a result of incremental costs due to control of pollutants, the
incremental capital investment, using utility financing is $15. 8 million;
using the discounted' cash flow (CSF) method, it is $16. 8 million. The
incremental net operating costs, after taking credit for sulfur recovered
at $10/long ton and phenol at 4c7lb, is $251, 800 (i. e., the pollutant
treatment facilities show a net annual operating profit).
* The costs presented here are preliminary estimates; the actual
costs of emission control in this process will not be known until
a full-scale facility has been constructed and operated.
** Incremental costs in this report refer to that portion of the total
plant cost that is assignable to pollution control facilities.
DC-20
-------
Table IX-8
Incremental Capital Investment3
($ Million)
Item
$x 106
Incremental plant investment
Wastewater treatment
Claus sulfur recovery with SCOT tail
gas treatment
Subtotal plant investment
Project contingencies0
(15% of the plant investment)
Incremental plant investment (I).
Startup costsd
(20% of incremental gross operating costs)
Interest during construction (IDC)e
Incremental working capital (W)'
Incremental capital investment (C)
3.0
8.4
11.4
1.7
13.1
0.3
Utility
13.1
0.3
DCF
3.1
0.3
16.8
Notes:
a Incremental plant investment, return on investment during con-
struction, and working capital are treated as capital costs in
year 0 (the year ending with completion of startup operation).
b Installed costs, including engineering design costs, contractors'
profit, overhead, and licences with no contingencies.
c Includes costs for unexpected site preparation and hardware
requirements at 15%. of plant investment.
d At 20% of incremental gross operating cost.
e For the DCF method, computed as the discount rate x incremental
Plant Investment for 1.875 years' average construction period. •
I(l-Ki)n = I(l.l2)'-875 = 1.236761 or 0.23676II additional investment.
f For the utility financing method, computed as the interest
rate on debt x incremental plant investment x 1.875 yrs.
Sum of materials and supplies at .9% of incremental plant invest-
ment and net receivables at 1/24 of annual incremental revenue.
IX-21
-------
Table EX-9
Incremental Annual Operating Costs ($)
Item
$/Year
Direct operating labor
(3 men/shift x $5/hr x 8, 304 shift hr/man-yr)
Maintenance labor (1.5% of I)
Supervision (15% of direct operating and
maintenance labor)
Incremental labor cost
Administration and general overhead
(60% of incremental labor)
Other direct costs
Supplies
Operating (30% of direct operating labor)
Maintenance (1.5% of I)
Incremental cost of supplies
Local taxes and insurance (2.7% of I)
Incremental gross operating cost
By-product credits
Sulfur (261.12 short tons/day x 365 days/yr x
0.9 capacity factor x $10/long ton x
2000/2240 long tons/short tons)
Phenol (72,000 Ib/day x 365 days/yr x
0.9 capacity factor x $0.04/lb)
Less by-product credit
Incremental gross operating cost
Incremental net operating cost (N)
124,600
171,000
44,400
37,400
171,000
765,900
946,100
340,000
204,000
400,000
208,400
307,800
1,460,200
-1,712,000
-251,800
IX-22
-------
The derivations of the formulae for calculating the incremental
cost of the syn-coal produced, due to pollution control, are presented
in Chapter III or this report. For the discounted cash flow method,
the required incremental annual cost of gas, X, for the assumed rate
of return is:
X = N + 0. 238161 + 0. 1275S 4 0. 230777W
where:
N = incremental net operating cost
I = incremental plant investment
S = startup costs
W = incremental working capital.
For the utility financing case, it is:
X = N + 0.1198C + 0. 0198W
where;
C = required incremental capital investment
N and W are as defined above.
Annual syn-coal production, G, for this analysis is:
(488, 376 Ib/hr) (15, 768 Btu/hr) (. 9 capacity
factor) (24 hrs/day) (365 days/yr) = 60.712 x
10!2 Btu/yr = 15. 300 x lO1^ kcal/yr
This can be regarded as a typical production rate for a
commercial SRC plant. Therefore, the incremental cost of the
product syn-coal, due to pollution control is:
annual cost of syn-coal _ X _ X
annual syn-coal production G 15. 3 x 10*2 kcal/yr
Solutions to the above equations for the incremental cost of
pollution control are shown in Table IX-10. Using the DCF method,
the resulting cost is 4. 9c7million Btu (19. 4c7million kcal); using
utility financing, it is 2. 7cVmillion Btu (10. 8c7 million kcal).
JX-23
-------
Table IX-10
Incremental Cost of Solvent Refined Coal
Product (SRC) Due to Pollution Control
Accounting Method
DCF
Utility financing
Incremental Cost of SRC Product
($/yr)
2,975,700
1,645,000
($/106 Btu)
0.049
0.027
The fuel gas evolved in the SRC process is consumed in the
process as feed for preheater boilers and steam generators. The
MEA system, therefore, should be considered as part of the pollution
control costs. In some process variations, however, this fuel gas
could be used for generation of makeup hydrogen, requiring the MEA
system to minimize catalyst poisoning. The relatively minor cost
of the MEA system has therefore been omitted from the control costing
in this analysis.
The cost of the Wellman-Lord and Allied Chemical processes
for independent treatment of the stack gas from the combustion of the
mineral residue to product steam were not considered either since
these alternatives were not selected by the process developer.
5. REFERENCE
"Economic Evaluation of a Process to Produce Ashless, Low-
Sulfur Fuel from Coal, " R&D Report No. 53, Interim Report
No. 1, Office of Coal Research.
IX-24
-------
X. THE TEXACO PARTIAL OXIDATION PROCESS
The Texaco partial oxidation process, also known as the Texaco
Synthesis Gas Generation process, is designed to convert high sulfur
residual oil into synthesis gas or hydrogen. It is licensed by the
Texaco Development Corporation, though a. competitive process is
licensed by the Shell Oil Company.
Development of the Texaco partial oxidation process began in
the late 1940's for noncatalytic conversion of methane to synthesis gas.
At that time the partial oxidation process was directly competitive with
methane-steam reforming, which utilizes catalysts packed in tubes that
are externally fired. The versatility of the partial oxidation process
has been expanded to include hydrocarbon feeds from methane to light
hydrocarbon gases, petroleum distillates, and residuum. Although a
solid fueled commercial plant is not yet in operation, Texaco states
that the process will accept lignite, bituminous, and anthracite coals.
The process has been accepted throughout the world. In the past,
the cost of synthesis gas generated by partial oxidation has been some-
what higher than that of gas produced by steam reforming of methane.
However, this situation is expected to change as a result of the avail-
ability of natural gas. The capability of partial oxidation plants to
accept high sulfur residual feedstocks represents an important ad-
vantage.
1. PROCESS DISPLAY
The Texaco partial oxidation process consists of three basic
sections:
The feed system, which delivers controlled amounts of
fuel, oxygen, steam and other materials to the generator
The refractory-lined reactor (gas generator), where fuel
reacts with steam and oxygen to produce hydrogen and
carbon monoxide in concentrations of 90 to 98 percent; its
capacity can vary between 14. 2 to 35. 4 x 10 m^/hr (0. 5
to 1. 25 x 10^ ft^/hr) of hydrogen equivalent product
X-l
-------
The cooling system that recovers sensible heat contained
in the reactor effluent gases.
(1) Bases for Analysis
The flow diagrams for two of the many possible variations
for the Texaco partial oxidation process, in synthesis gas and
hydrogen production modes, are presented in Figures X-l and
X-2, respectively, along with stream compositions. Note, how-
ever, that either process mode can be adapted readily to the pro-
duction of the other product. The synthesis gas and hydrogen
production modes shown in Figures X-l and X-2 use the same
quantities of feedstock; both process variations were included
to indicate process versatility. The respective sulfur and energy
balances are shown in Tables X-l and X-2. The feedstock is
assumed to be a 6.2 percent sulfur residual fuel oil. The sulfur
flow rate is indicated in brackets on the flow diagram for selected
streams.
For the producing of synthesis gas, 98. 4 percent of the
total sulfur in the feed oil is recovered in its elemental form for
sale; the remaining 1. 6 percent is released to the atmosphere.
For the hydrogen generation process mode, 98 percent of the
sulfur is recovered for sale and 2 percent is vented to the atmo-
sphere.
The process is discussed in more detail below under the
following headings:
Process description - synthesis gas
Process description - hydrogen
Sulfur and energy balances
Pollution control processes
Cost of pollution control.
(2) Layout and Symbols
The general direction of the process flow is from the
residual feed oil on the left to the cleaned product fuel on the
right. This primary conversion stream is indicated by a bold
line. All residuals and by-products are shown along the bottom
of the flow sheets (inverted trapezoids).
X-2
-------
FIGURE X-1
The Texaco Process
Synthesis Gas Production
COMPOSITION OF GASEOUS STREAMS (VOL %)
COMPONENT
H2
CO
co,
CH4
COS
^
HZS
H,0
NH,
°2
Nj + A
FLOW RATE
m3/hr X 1()P
ft3/hr X 10^
0
39.49
44.94
4.34
0.56 .
0.07
-
1.30
9.53
28pp>n
-
0.17
1.101
38.89
©
43.15
48.66
4.74
0.61
0.08
-
1.42
1.15
-
-
0.19
1.003
36.60
©
45.79
61.65
0.05
0.65
50 pom
-
10 pom
1.66
-
-
0.20
0.950
33.54
©
46.52
52.46
0.05
0.66
-
-
-
aii
-
-
0.20
0.935
33.02
©
-
-
11.72
-
-
0.02
-
11.02
-
2.59
74.65
0.148
5.21
©
-
-
69.70
-
-
-
21.84
8.46
-
-
-
0.069
2.44
COMPOS. OF LIOUIDSTREAMS IWT *>
COMPONENT
C
H
S
N
Am
FLOW RATE
Ihort ton/nr
fnton/hr
A A
83.63 84.00
10.02 15.97
6.20 0.03
0.23
0.02
364.92 9.06
331.06 S.20
COMPONENT
HC
COO
TOS
HjS
NH,
Ash
FLOW RATE
thort ton/hr
— *"
A
lOiom
Sppm
0.5
25ppm
0.03
0,1 ppm
80.38
7Z92
Non: [ I irKfieata tulfur ftowrne, in m torn per nour
LEGEND
n?
S
TDS
COO
Cubic menn
Sulfur
Totti diodvid aunt
Chemical oxygen demand
-------
FIGURE X-2
The Texaco Process
Hydrogen Production
_J
COMPOSITION OF GASEOUS FUELS (VOL *>
COMPON.
- COMPONENT -
H2
CO
CO,
CH4
COS
so,
H,S.
H,0
NH3
°2
Nj + A
FLOW RATE '
m3/hr X 10*
H3/hi X 108
©
39.49
44.54
4.34
0.56
0.07
-
1.30
9.53
28ppm
-
0.17
1.101
38.89
©
23.84
26.88
2.62
0.34
0.04
-
0.79
45.39
-
0.10
1.825
64.43
©
61.81
0.95
35.56
0.42
0.05
-
0.97
0.11
-
-
0.13
1.475
52.07
©
61.42
0.94
34.08
0.41
0.01
-
3pfxn
-
-
-
3.14
1.484
52.41
© ©
97.69
1.50
0.01 12.36
0.60
-
0.042
Ippm
9.90
-
2.60
0.20 75.10
0.932 0.894
32.91 31.56
© ©
0.20
54.42 91.58
0.10
2.00 0.02
-
42.05 5 nxn
-
-
-
I.U 8.10
0.034 "0.552
1.20 19.50
COMPOS. OF LIQUID STREAMS IWT %>
COMPONENT
C
H
s
N
Adi
FLOW RATE
A A
83.53 84.00
10.02 15.97
6.20 0.03
0.23
0.02
*K»iton,1u : 364.92 63.2
mwrt/hr
COMPONENT
HC
COO
TDS
V
NH3
Atfl
FLO* RATE
(hort ton/hr.
A
10pptn
5ppm
0.5
25P.X,
0.03
0.1 ppm
80.38
331.06 57.341 mion/hr 72.92
Not*: [ I indfcatts tulfur How rat* in m ton* per hour
LEGEND
m cofw
TDS
COO
Metric ton,
Total dissolved solids
Chemical oxvgen demand
-------
Table X-l
Sulfur Balance, Short Tons/Hr (Metric Tons/Hr)
Carrier
Input
Residual fuel oil
Low-sulfur oil
Output
Elemental sulfur
Sulfur dioxide to atmosphere
Sulfur compunds from Clans tail
gas to atmosphere
Sulfides from iron sponge drums
vented to atmos. (primarily 802)
Synthesis Gas
22.593(20.496)
22.59(20.493)
.003 (.003)
22.593 (20.496)
22.24(20,178)
0.053 (0.048)
0.22 (0.20)
, 0.08 (0.07)
Hydrogen
22.609(20.51)
22.59 (20.493)
0.019 (.017)
22.609(20.51)
22.15(20.09)
0.069 (.063)
0.220 (.200)
0.1 70 (.154)
The circled figures refer to the gaseous streams while
triangles refer to the liquid streams. Their compositions are
tabulated beneath each flow sheet. The rhombic-shaped units
represent intermediate products and sources for which the
distribution is not shown. Nonintegral pollutant cleanup pro-
cesses are indicated by sloping rectangles. The dashed line
encloses that portion of the process defined by the process de-
signer. The remaining unit processes were assumed for this
analysis.
2. PROCESS DESCRIPTION - SYNTHESIS GAS
This section describes the specific steps of the Texaco partial
oxidation process in a synthesis gas production mode. The pollution
control and cleanup processes are similar for the two process modes
and are discussed jointly in a subsequent section.
X-5
-------
Table X-2
Energy Balance *
Carrier
Input
Residual fuel oil
Low-sulfur oil
Output
Synthesis gas
Hydrogen
Sulfur
Ammonia
Cooling water
Steam
Other (by difference)*
Heating Value
10.000 kcal/kg
1 0.560 kcal/kg
2,903 kcal/m3
2.913 kcal/m3
2, 213 kcal/m3
5,375 kcal/kg
19.5 kcal/1
647 kcal/kg
—
SYNTHESIS GAS
Mass
Flow Rate
331.060 kg/hr
8.210kg/hr
0.935 x 106m3/hr
—
20 ,1 76 kg/hr
22 kg/hr
11 x 106I/hr
—
—
Energy Flow Rate
(kcal/hr x 106)
3.397
3.310
87
3.397
2.714
—
45
.118
213
—
424.88
(Btu/hrx 106)
13.480
13,134 '
346
13.480
10.770
—
178
0.467
847
—
1684.5
HYDROGEN
Mass
Flow Rate
33 1 ,060 kg/hr
5 7.330 kg/hr
0.932 x 106m3/hr
20,090 kg/hr
22 kg/hr
1 1 x 106 I/hr
2 13, 190 kg/hr
—
Energy Flow Rate
(kcal/hr x 106)
3,916
3,310
606
3,916
2,715
44
.118
213
138
805.88
(Btu/hr x 106)
15,539
13,134
2,405
15,539
10,773
175
0.467
847
547
3196.5
t For Synthesis gas mode, the overall efficiency is 81.2% and 70.5% for the hydrogen mode.
"Includes heating values of other products, sensible heat of product streams, and heat lost directly to the atmosphere.
-------
(1) Generation
Most of the feed oil, plus soot-laden oil returned from the
scrubber located downstream of the generators, enters the gen-
erator burners, along with oxygen and steam; a portion of the
feed oil is directed to the downstream scrubber. All of these
inputs are preheated to minimize oxygen consumption. The
oxygen and feed oil streams are not premixed to avoid problems
with flow distribution, flash-backs, preheat, and pressure limits.
Pressures in the generator may range from near-atmospheric to
86kg/cm^ (1200 psig), while maximum temperatures are limited
by steam injection to between 1100 to 1600 degrees C (2000 to
2800 degrees F).
The hot raw output gas contains mainly carbon monoxide
and hydrogen and small amounts of carbon dioxide, as well as
methane, carbonyl sulfide, hydrogen sulfide, nitrogen, argon,
and soot impurities.
(2) Scrubbing and Soot Recovery
The hot gas stream from the generator is cooled by passage
through a waste heat recovery (WHR) section, where it raises
steam to drive the air separations plant. The gas is then scrubbed
with a portion of the feed oil to remove the soot, before continuing
on to the purification stage.
The soot-laden oil is separated in the soot recovery unit
into two fractions to maintain a steady-state removal of ash from
the system. * This process mode has employed oil-washing of
the soot (although the washing mode has not been commercially
utilized). Texaco uses water-washing as presented in the alter-
nate production mode. The smaller, soot-rich fraction serves
as fuel for the steam plant, while the cleaner fraction is returned
to the gas generator for further processing.
(3) Purification
The purification stage that follows is proposed for purposes
of this analysis, and is not a part of the Texaco process as
The process licensor claims that this flow scheme is not neces-
sary to eliminate ash from the system.
X-7
-------
conceived by the developers. A hot carbonate process can be
used to treat the high H2S/CO2 ratio in this feed gas and also
minimize COS emissions, The gas stream from the oil scrubbers
flows through a low temperature waste heat recovery section to
condense out moisture and to preheat boiler feed water for the
steam plant. The gas is then scrubbed at 93 to 105 degrees C
(200 to 220 degrees F) in a hot potassium carbonate absorber to
remove carbon dioxide and hydrogen sulfide and to hydrolyze much
of the carbonyl sulfide present. The scrubbed gas is again cooled
to condense out any remaining moisture, and then passed through
two iron sponge beds, to remove final traces of sulfur.
The hot potassium carbonate solution is regenerated with
low pressure steam and the regenerator effluent gas is directed
to the sulfur recovery section. The sulfur retained by the iron
sponge beds is vented to the atmosphere during periodic regen-
eration. The synthesis gas product is free of sulfur and con-
tains approximately 52. 5 percent carbon monoxide and 46. 5 per-
cent hydrogen.
(4) Steam and Oxygen Generation
The overall steam requirements of the process—generator
steam and G£ separation — can be met, in part, by the first waste
heat boiler. An additional oil-fired steam generator would sup-
ply the remainder. Fuel for the steam generator consists of the
soot-laden oil from the soot recovery section, with enough low-
sulfur fuel oil added to maintain stack gas emissions at 0.5 Ib
SO2/million Btu of fuel burned or 0.9 kg/10^ kcal (equivalent to
0.4 percent sulfur in the fuel oil).
An air separation plant powered by steam from the first
waste heat recovery section serves as a source of oxygen for
the synthesis gas generator.
3. PROCESS DESCRIPTION - HYDROGEN
This section describes the Texaco partial oxidation process in
the hydrogen production mode. Some of the step descriptions are
similar to those for synthesis gas production and are not repeated
here.
X-8
-------
(1) Generation
This generation step essentially duplicates that for synthesis
gas production. The major distinction lies in the provision of a
water scrubber in the lower portion of the generator vessel.
(2) Scrubbing and Soot Recovery
The hot gas stream from the gas generator is scrubbed in
two consecutive water scrubbers, where it picks up sufficient
moisture to feed the subsequent CO-shift reaction.
The soot-^laden slurry from the first scrubber, located in
the lower portion of the generator vessel, flows to the soot re-
covery unit, where it is scrubbed by naphtha to extract the soot.
The water is then directed to the wastewater treatment unit and
eventually recycled to the second scrubber. The soot-laden
naphtha is mixed with feed oil and distilled to yield naphtha and
soot-laden feed oil separately. The latter is separated into two
fractions, as in the synthesis gas production mode. The soot-
rich fraction serves as fuel for the steam plant, while the re-
maining fraction is returned to the gas generator for further
processing. This soot recovery scheme is alternative to the
oil washing presented in the other mode.
Water for the second scrubber is fed from the wastewater
treatment stage and condensate from the waste heat recovery
boiler on the 'CO-shift effluent. Its wastewater serves as the
water feed for the first scrubber.
(3) CO-Shift Conversion
The shift conversion step increases the H2/CO ratio from
0. 9 to 65. It takes place with the aid of a cobalt-molybdenum
catalyst, which is more effective than conventional iron oxide
catalysts in the presence of sulfur. The effluent gas stream is
cooled by passage through a waste heat recovery (WHR) unit that
generates low pressure steam. The moisture condenses out and
serves as partial feed for the second scrubber.
X-9
-------
(4) Purification
In this presentation of the Texaco process, removal of car-
bon dioxide and hydrogen sulfide impurities can be achieved by
means of two Rectisol process stages. The product hydrogen
stream contains 97 to 98 percent hydrogen, 1 to 2 percent carbon
monoxide, and only very minor carbon dioxide and hydrogen
sulfide impurities. The Rectisol system consists of a refrig-
eration section that cools the gas to -20 to -60 degrees C
(0 to -80 degrees F), a methanol absorber and a section that
regenerates the methanol.
4. POLLUTION CONTROL PROCESSES
The nature and treatment of major waste streams is presented
in Table X-3 and discussed briefly below. A more complete discussion
of the specific processes introduced here is given in Chapter III.
Table X-3
Nature and Treatment of Major Waste Streams
Waste
Hydrogen sulfide
Sulfur dioxide
Carbonyl sulfide
Wastewater
Ash
Solid sludge
Iron sponge
CO shift catalyst
Sources
Regeneration of acid-gas absorbers
Wastewater treatment
Beavon tail gas
Steam plant
Regeneration of iron sponge
Regeneration of Rectisol II
Soot recovery
Gas condensates
Boiler and cooling tower blowdown
Steam plant
Wastewater treatment
Iron sponge drum
Shift converter
Treatment
Claus and Beavon
Claus and Beavon
None required
None required
None required
None required
Chevron and biological
None required
Oxidation
Oxidation
Oxidation
X-10
-------
(1) Gaseous Wastes
The hydrogen sulfide in the effluent streams from the acid^
gas removal processes is converted to liquid sulfur in the sulfur
recovery section consisting of a Glaus plant and a Beavon unit
to process the Claus tail gas, before it is vented to the atmo-
sphere. Approximately 99 percent of the incoming sulfur is
recovered.
Sulfur dioxide emissions from combustion of the mixture
of soot-laden feed oil and low sulfur fuel oil are considered too
small to require treatment under current standards. ,The
effluent stream of the Rectisol II stage also contains low con-
centrations of sulfur; however, much of that sulfur is in the
reduced form. Incineration of this stream would probably
require excessive fuel consumption due to the quantity of
present.
(2) Liquid Wastes
Wastewater from the soot recovery unit in the hydrogen
production mode of the process, as well as gas condensates and
blowdown water from steam generators and cooling towers,
are treated in the wastewater treatment plant. The plant includes
a Chevron unit and a biological treatment unit and yields an inert
solid residue and small amounts of ammonia and hydrogen sulfide
gases. The latter is treated by the sulfur recovery section.
(3) Solid Wastes
Because the feedstock is residual fuel oil, generation of
ash is minimal. The ash is eventually rejected by the steam
plant and can be used as landfill or sold to metal processors for
recovery of vanadium and nickel. The solid sludge from the
wastewater treatment plant requires oxidation prior to landfilling.
The spent CO shift catalyst and iron sponge should be fully oxi-
dized prior to removal to neutralize their pyrophoric properties.
X-ll
-------
5. COSTS OF POLLUTION CONTROL*
The costs of the pollution control and treatment processes de-
scribed in Sections 2, 3 and 4 of this synopsis are calculated here
using both a utility financing and a discounted cash flow accounting
method (DCF). The major assumptions and conventions adopted here
are discussed in Chapter III of this report. All by-product flow rate
data presented in Figures X-l and X-2 also apply.
The incremental** capital and operating costs for pollution con-
trols for the Texaco partial oxidation process are reported for the
two cases considered above: production of synthesis gas and produc-
tion of hydrogen.
Capital investment requirements are shown in Tables X-4 and
X-5, Using the utility financing method, the capital investment re-
quired for pollution control is $16. 0 million and $12. 7 million, for
producing synthesis gas and hydrogen respectively. Using the DCF
method, it is $17. 1 million and $13.4 million. Tables X-6 and X-7
present calculations of the annual operating costs for both product
streams. For synthesis gas production, an additional $22,300 per
year is required for pollution control while a net savings of $172,000
is expected when producing hydrogen.
The derivations of the formulae for calculating the incremental
cost of gas production due to pollution control are presented in Chap-
ter III of this report. For the DCF method, the required incremental
annual cost of gas, X, for the assumed rate of return is:
X = N + 0.238161 + 0. 1275S + 0. 230777W
* The costs presented here are preliminary estimates; the actual
costs of emission control in this process will not be known until
a full-scale facility has been constructed and operated.
>:<>:< Incremental costs in this report refer to that portion of the total
plant cost that is assignable to pollution control facilities.
X-12
-------
Table X-4
— o
Incremental Capital Investment ($Million)
For Synthesis Gas Production
(For 6. 2% Sulfur in Residual Oil Feedstock)
Item
Incremental plant investment
Wastewater treatment
Claus sulfur recovery with
tail gas treatment
Project contingencies0
(1 5% of the plant investment)
Incremental plant investment (I)
Startup costs
(20% of incremental gross
operating cost)
Interest during construction2 (IDC)
Incremental working capital (W)
Incremental capital investment (C)
$ Million
3.0
8.6
11.6
1.7
13.3
0.3
Utility
2.2
.2
16.0
13.3
0.3
DCF
3.2
.3
17.1
Notes:
a Incremental plant investment, return on investment during construction,
and working capital are treated as capital costs in year 0 (the year ending
with completion of startup operation).
b Installed costs, including engineering design costs, contractors' profit,
overhead, and licenses with no contingencies.
c Includes costs for unexpected site preparation and hardware requirements
at 15% of plant investment.
d At 20% of incremental gross operating cost.
e For the DCF method, computed as the discount rate x incremental
plant investment for 1.875 years' average construction period.
I(l+i)n = 1.12)1-875 = 1.236761 additional investment.
For the utility financing method, computed as the interest rate
on debt x incremental plant investment x 1.875 yrs.
f Sum of materials and supplies at .9% of incremental plant investment
and net receivables at 1/24 of annual incremental revenue.
X-13
-------
Table X-5
Incremental Capital Investmenta ($ Million)
For Hydrogen Production
(For 6. 2% Sulfur in Residual Oil Feedstock)
Item
$ Million
Incremental plant investment"
Wastewater treatment
Glaus sulfur recovery with
tail gas treatment
Subtotal plant investment
Project contingencies0
(15% of the plant investment)
Incremental plant investment (I)
Startup costsd (S)
(20% of incremental gross operating cost)
Interest during construction6 (IDC)
Incremental working capital' (W)
Incremental capital investment (C)
3.0
6.0
9.0
1.4
10.4
0.3
Utility
1.8
.2
12.7
10.4
0.3
DCF
2.5
.2
13.4
Notes:
a Incremental plant investment, return on investment during construction,
and working capital are treated as capital costs in year 0 (the year ending
with completion of startup operation).
b Installed costs, including engineering design costs, contractors' profit,
overhead, and licenses with no contingencies.
c Includes costs for unexpected site preparation and hardware requirements
at 15% of plant investment.
d At 20% of incremental gross operating cost.
e For the DCF method, computed as the discount rate x incremental plant
investment for 1.875 years' average construction period.
I(l+i)n = I(1.12)L875 = 1.236761 additional investment.
f. Sum of materials and supplies at .9% of incremental plant investment and
net receivables at 1/24 of annual incremental revenue.
X-14
-------
Table X-6
Incremental Annual Operating Costs of
'•' Synthesis Gas Production
Item
$/Year
Direct operating labor (3 men/shift x
$5/hr x 8,304 shift hr/man-year)
Maintenance labor (1.5% of I)
Supervision (15% of indirect operating
and maintenance labor)
Incremental labor cost
Administration and general overhead
(60% of incremental labor)
Other direct costs*
Supplies
Operating (30% of direct operating labor)
Maintenance (1.5% of I)
Incremental cost of supplies
Local taxes and insurance (2.7% of I)
Additional cost of fuel by using naphtha
rather than resid (64.6 bbl x $3 diff/bbl)
Incremental gross operating cost
By-product credits
Sulfur (533.76 short tons/day x 365 days/year x
.9 capacity factor x $10/long ton x 2000/2400
long tons/short tons)
Ammonia (.576 short tons/day x 365 days/year x
.9 capacity factor x $25/short ton)
Total by-product credit
Incremental gross operating cost
Less by-product credits
Incremental net operating cost (N)
124,600
199,500
48,600
37,400
199,500
372,700
223,600
400,000
236,900
359,100
200
1,592,500
1,565,500
4,700
1,570,200
1,592,500
-1,570,200
22,300
This item includes incremental costs of chemicals, catalysts and make-up water
X-15
-------
Table X-7
Incremental Annual Operating Costs of
Hydrogen Production
Item
$/Year
Direct operating labor (3 men/shift x
$5/hr x 8,304 shift hr/man-yr)
Maintenance labor (1.5% of I)
Supervision (15% of indirect operating
and maintenance labor)
Incremental Labor Cost
Administration and general overhead
(60% of incremental, labor)
Other direct costs*
Supplies
Operating (30% of direct operating labor)
Maintenance (1.5% of I)
Incremental cost of supplies
Local taxes and insurance (2.7% of I)
Additional cost of fuel by using naphtha rather
than resid (451 bbl x $3 diff/bbl)
Incremental gross operating cost
By-product credits
Sulfur (531.6 short tons/day x 365 days/
yr x .9 capacity factor x $10/long ton x
2000/2240 long tons/short tons)
Ammonia (.576 short tons/day x 365 days/
yr x .9 capacity factor x $25/short ton)
Total by-product credit
Incremental gross operating cost
Less by-product credits
Incremental net operating cost (N)
124,600
156,000
42,100
37,400
156,000
322,700
193,600
400,000
193,400
280,800
1,400
1,391,900
1,559,200
4,700
1,563,900
1,391,900
-1,563,900
-172,000**
*This item includes incremental costs of chemicals, catalysts and make-up water
**By-product credits exceed operating costs and result in a net annual operating profit
X-16
-------
where
N = incremental net operating cost
I = incremental plant investment
S = startup costs
W = incremental working capital.
For the utility financing case, the appropriate formula is:
X = N + 0. 1198C + 0.0198W
where
C = required incremental capital investment
N and W are as defined above.
Solutions to the above equations for the incremental cost of gas are
shown on Tables X-8 and X-9.
For the DCF method, the incremental cost of gas is 3. 9c7l06 Btu
for synthesis gas production and 2. ScVlO^ Btu for hydrogen production.
For the utility financing case, these costs are 2. 3
-------
Table X-8
Incremental Cost of Synthesis Gas Production
Due to Pollution Control
(6.2% Sulfur in Residual Oil Feedstock)
Accounting Method
DCF
Utility financing
; Cost of Gas
($/yr)
3,297,000
1,943,100
($/106Btu)
.039
.023
Table X-9
Incremental Cost of Hydrogen Production
Due to Pollution Control
(6. 2% Sulfur in Residual Oil Feedstock)
Accounting Method
DCF
Utility financing
Cost of Gas
' ($/yr) '
2,389,400
1,353,500
($/106 Btu)
.028
.016
X-18
/•..
-------
XI. THE DESULFURIZATION OF CRUDE OIL
Until about 1950, desulfurization in petroleum refining was
directed primarily at the reduction of sulfur in light products such as
gasoline, naphtha, and kerosene. In the case of gasoline, desulfuri-
zation improved the product odor and achieved better tetraethyl lead
susceptibility, therefore reducing the cost of producing gasoline and
increasing customer acceptance. Some chemical treatment methods,
such as sulfuric acid washing, improved product appearance and
stability and also reduced sulfur content to some extent. At that time,
the sulfur compounds of primary concern were the foul-smelling
mercaptans (compounds analogous to alcohols and phenols which contain
sulfur in place of oxygen). Several processes were used to convert
these mercaptans to disulfides which had a more acceptable odor.
Although the total sulfur content of the oil remained the same, it was
now more marketable. Some selective solution processes were also
developed to remove mercaptans. During this period, only the simplest
forms of catalytic desulfurization were used. A typical catalytic
process would contact vaporized oil with a fixed bed of bauxite. Part
of the sulfur was converted to H^S which was removed by caustic
washing or distillation.
In 1948, the first commercial platinum catalyst reforming
process was put into commercial operation. This was the Platforming
process, developed by the Universal Oil Products Co. By 1950,
catalytic reforming was accepted throughout the refining industry and
was being rapidly installed.
The purpose of catalytic reforming was to increase the octane
rating of gasoline and naptha. The main reactions occurring in this
process are dehydrogenation of aromatics (hydrocarbons with at least
one benzene ring, CgHL), dehydrocyclization of paraffins to form
aromatics (breaking down heavier hydrocarbons to aromatics), iso-
merization (formation of isomers) and hydro era eking.. As a net
result of these reactions, an excess of hydrogen-rich gas is formed
which must be vented from the process. When catalytic reforming
is installed only for increasing octane number, the hydrogen-rich
off-gas is produced at fuel gas value. With the availability of cheap
hydrogen from catalytic reforming came the first stirring of wide
interest in hydrodesulfurization (HDS) processing.
XI-1
-------
In some refineries, this excess hydrogen was used in processes
such as the Hydrobon process. Feedstocks intended for catalytic re-
forming were first processed through a Hydrobon unit to reduce the
sulfur content to low levels and also to reduce the content of nitrogen
compounds. Both sulfur and nitrogen acted as poisons to the catalytic
reforming catalyst. In addition to catalytic reformer feedstock treat-
ment, excess hydrogen was used to upgrade kerosene and jet fuel,
lubricating oils, waxes, fuel oils, and most recently, residual stocks.
As the scope of hydrotreatment expanded, other sources of hydrogen
were needed to supplement hydrogen from reforming.
1. PROCESS DISPLAYS
Five process flow diagrams for different desulfurization
schemes are presented in Figure XI-1 and discussed in Section 2. The
desulfurized oil's sulfur content decreases, in general, as process
complexity and costs increase. The zone contained within the dashed
lines represents the desulfurization processes considered in this study.
The plant sulfur balance for each scheme is also presented on the
process flow diagrams.
(1) Bases for Analysis
It should be emphasized that no two refineries are
identical. The flow rates and compositions presented in this
study are for a hypothetical plant. They are the result of
a particular combination of feedstock, processing operations
and product mix selected for this analysis. Care must be
exercised and allowances made when attempting to use this
data for variant applications.
Assumptions made in preparing the desulfurization flow
sheets include:
Crude oil feed rate of 100, 000 bbl/day. This size
was selected since present refinery economics
dictate large throughputs; the other processes
discussed in this study reach a cost plateau at
lower energy input rates. This feed rate, on an
energy basis, is about 50 percent higher than the
coal-based processes discussed in other clean-fuel
evaluations. Energy output rates are also about
doubled. The costs presented in Section 4 are.
XI-2
-------
FIGURE XI-1, Case 1
Desulfurization of Crude Oil
U)
LEGEND
S SULFUR
BCD BARRELS PER DAY
( I SULFUR FLOW RATE
IN METRIC TONS PER DAY
FOE FUEL OIL EQUIVALENT
— k-
S*
1 '
^
TOPPING
UNIT
L
i
,*•
(thai
LJ
^
[3.8]
Gel Oil
• I7fl7l
•Wii
[199.11
^
ATM
Residuum
[0.9]
J
21,100 BPD
?3^Pr
1%S
*"
H2
\. PLANT ./^~
^/
93.450m3/Dav
Ln
\ GASOLINE /
20.300 BCD
64° API. 0.04% S
S \
HYDflOTREATER
L^ ^)
s^" ~^\
HDS
. .
\ '
44.000 BPD
17.5° API
3XS
/ OA
1 SUPF
/
[0.41
t
f
[3.4]
[3.11
^•^
/
[256)
I
V u
GAS
MEA
S~^\
.c
U
B
G
(29) ASH2
"
V
I* \
46,725m3/Div
K002I
\ FUEL GAS /
\ SULFUR
I 9B FOE BARRELS/DAY 128.7]
[0.3] AS
so,
1
SULFUR
]
I
DESULFURIZEO\
OIL \
1.66Wl%S
1
PLANT SULFUR BALANCE, METRIC TONS/DAY
IN
100,000 BPD CRUDE, 1,7% S, 32.6° API 232.5
SULFUR OUT
GASOLINE (20,300 BPD, 0.04% S, 64° API) OS
DESULFURIZEDOIL
TOPPING UNIT BOTTOMS 199.1
j V FROM HDS UNIT 3.1 202.6
/ / A™ RECOVERED FROM FUEL GAS 28.7
/ / RELEASED TO ATMOSPHERE O.3
J
TOTAL 232.5
-------
FIGURE XI-1, Case 2
Desulfurization of Crude Oil
X
S SULFUR
BPD BARRELS PER DAY
f I SULFUR FLOW RATE
IN METRIC TONS PER DAY
FOE FUEL OIL EQUIVALENT
DESULFUR1ZED >
OIL
POOL
PLANT SULFUR BALANCE. METRIC TONS/DAY
20.300 BPD
64° API. 0.04% S
100.000 BPD CRUDE. 1.7% S. 32.6° API
SULFUR OUT
GASOLINE 120.300 BPD. 0.04% S. 64° API)
OESULFURIZEDOIL
VACUUM STILL BOTTOMS
FROM HYDROTREATER
FROM HDS UNIT
RECOVERED FROM FUEL GAS
REMAINING IN FUEL GAS
RELEASED TO ATMOSPHERE
TOTAL
0.2
1.2
-------
FIGURE XI-1, Case 3
Desulfurization of Crude Oil
X
r
is*
5 SULFUR
BPD BARRELS PER DAV
f ) SULFUR FLOW RATE
IN METRIC TONS PER DAY
FOE FUEL OIL EQUIVALENT
PLANT SULFUR BALANCE, METRIC TONS/DAY
SULf UR IN
100,000 BPD CRUDE, 1.7%S, 32.6° API
SULFUROUT
. GASOLINE (20,300 BPD, 0.04% S, 64° API
OESULFURI2EOOIL
SOLVENT REFINING BOTTOMS
FROM HYDROTREATER
FROM HDS UNIT I
FROM HDS UNIT M
RECOVERED FROM FUEL GAS
REMAINING IN FUEL GAS
RELEASED TO ATMOSPHERE
33.9
0.4
12.4
176.1
0.2
TOTAL
307 FOE BARRELS/DAY
-------
FIGURE XI-1, Case 3A
Desulfurization of Crude Oil
X
*r
a\
LEGEND
S
BPD
( I
FOE
SULFUR
BARRELS PER DAY
SULFUR FLOW RATE
IN METRIC TONS PER DAY
FUEL OIL EQUIVALENT
00 BPD, 44.3° APt
I 0.22% S
1
HYDROTREATER
MEA ABSORBER
' AND
OESULFURIZEO >
OIL
POOL
I
PLANT SULFUR BALANCE. METRIC TONS'DAY
100.000 BPD CRUDE. 1.7* S. 32.6° API
SULFUR OUT
GASOLINE 120.300 BPD. 0.04% 5.64° API
DESULFURtZEOOIL
FROM HYDROTREATER
FROM HDS I
FROM HDS II
RECOVERED FROM FUEL GAS
REMAINING IN FUEL GAS
RELEASED TO ATMOSPHERE
TOTAL
0.4
12.4
19.6
209.7
0.2
2.1
307 FOE BARRELS/DAY
-------
FIGURE XI-1, Case 4
Desulfurization of Crude Oil
•X
S SULFUR
BPD BARRELS PER DAY
[ I SULFUR FLOW RATE
IN METRIC TONS PER DAY
FOE FUEL OIL EQUIVALENT
20,300 BPD
64° API. 0.04% S
OTREATEH
. J
\
>.
HDS
'
•"^
SJOUUM
IRECT
SULFUR 1-
ATI ON
[3.4] .
13.1]
/
[256]
[18.21
— *
\ ^
[180.9]
^
I NATURAL \
-/ GA
/ SUPP
/
S \
LY \
\
934.550m3/Day
MEA
r s 1
c
R
U
r
[209.4] AS
-
' .
K0.5]
\ FUEL GAS /
\ SULFUR
\ 1207.3]
828 FOE BARRELS/DAY
[2.1] AS
so2
1
OESULFURIZEO\
OIL \
POOL \
)
0.2* S
1
PLANT SULFUR BALANCE. METRIC TONS/DAY
SULFUR IN
SULFUR OUT
GASOLINE 120.300 BPD. 0.04* S. 64" API) 0.9
DESU
LFURIZEDOIL
DIRECT DESULFUHIZAT1ON 18.2
FROM HYOROTREATER 0.4
/ / ,TM RECOVERED FROM FUEL GAS 207.3
/ / REMAINING IN FUEL GAS 0.5
TOTAL 232.5
-------
therefore, not strictly comparable to the other pro-
cess evaluations when viewed on a unit-to-unit basis.
A crude oil containing 1. 7 percent by weight of sulfur,
gravity of 32. 6 degrees API was selected. This type
of crude oil is representative of a moderately high
sulfur content crude of a type which may be obtained
from the Persian Gulf or from many wells in the
United States.
It was assumed that the desulfurization was incre-
mental to an existing refinery which had excess
crude oil processing capacity. Therefore, the prob-
lems of wastes which are associated with the main
cooling water system, steam boiler system, or the
main wastewater treatment system have not been
considered.
(2) Layout and Symbols
The general direction of process flow is from the crude
oil feed on the left to the desulfurized oil on the right, with
residuals shown along the bottom (denoted by inverted trape-
zoids). The bold lines indicate the primary oil flow. Rhombic-
shaped units represent intermediate products, uses, and sources,
for which the distribution is not shown. Nonintegral pollutant
cleanup processes are indicated by sloping rectangles.
2. PROCESS DESCRIPTION
The crude oil is desalted (removal of soluble sodium, calcium
and magnesium salts) and fractionated in an atmospheric still where
the most volatile fraction, gasoline, is extracted. The 149 degrees C
(300 degrees F) and lighter material bypasses the desulfurization
process and is blended into the refinery gasoline pool. The gasoline
contains only 0. 04 percent sulfur. The remaining 79. 7 percent of
the crude oil is separated into various fractions and hydrodesulfurized.
The desulfurized oil products from these fractions are combined
XI-8
-------
with the residuum and identified as "desulfurized oil." This is not
a specification fuel oil product. The desulfurized oil represents a low
sulfur oil pool from which a variety of low sulfur distillates or fuel
oils may be derived by distillation or blending. Products separated
from the low sulfur oil may include gasoline, which was formed in
small amounts during hydrodesulfurization, kerosene and jet fuel,
light fuel oil containing very low sulfur, and higher sulfur heavy fuel
oil which might be blended with residuum. No attempt has been made
to identify the specific products and rates of production since these
change with refinery location, markets, and time.
Table XI-1 contains a listing of currently available hydrodesulfur-
ization processes. Different processes are applied to different boiling
ranges of oils. Lower boiling oils normally contain less sulfur than
higher boiling materials from the same crude oil. The sulfur in the
lighter oils exists in the forms that are more easily treated. The lower
boiling oils can be desulfurized more easily at lower pressures, lower
temperatures, and at higher catalyst space velocities (volume of oil
per hour per volume of catalyst). The most difficult hydrodesulfuri-
zation feedstocks are residual: the oils which cannot be distilled in
atmospheric or vacuum towers. Vacuum residuum usually has a
boiling range of 538°C (1000°F) and higher. Residual hydrodesulfuri-
zation is more difficult than distillate desulfurization because the
stocks are dirtier and the molecular structure is more complex. Non-
distillable impurities are contained in crude oil concentrate in the resid-
uum and must be handled by the hydrodesulfurization process. Unfor-
tunately, these impurities tend to plug the catalyst beds and may irre-
versibly poison the catalyst. The heavy residuum oil tends to deposit
carbon on the catalyst, reducing its activity. Finally, the residuum
stocks often have a higher demand for hydrogen than the distillates,
and the cost of hydrogen is of major importance for good hydrodesulfur-
ization economies.
Normal sources of hydrogen for petroleum refineries have been
by-product hydrogen from catalytic reforming or by-product hydrogen
from petrochemical processes such as ethylene-cracking. More
recently, steam reforming units for natural gas or refinery gas have
been installed to manufacture the hydrogen. In refineries where an
excess of naphtha may be available (mainly foreign refineries), .naphtha-
steam reforming has been used to manufacture hydrogen. The Texaco
and Shell partial oxidation processes can use almost any type of liquid
hydrocarbon feed. The versatility of partial oxidation insures that it
will be more widely used in the future.
XI-9
-------
Table XI-1
Important Hydrodesulfurization Processes
i
*_*
o
Licensor
Chevron Research Co.
Cities Service R&D Co.
and Hydrocarbon Research Inc.
ESSO Research & Engineering Co.
and Union Oil of California
Gulf Research & Development Co.
Institute Francais de Petrole
Standard Oil Co. (Indiana)
Universal Oil Products Co.
Process
VGO Isomax
RDS Isomax
VRDS Isomax
H-Oil
Go-Fining
Residfining
Gulf HDS Type I and Type II
Gulf HDS Type III
Heavy Gas Oil Gulfining
IFP Vacuum Gas Oil HDS
IFP Resid HDS
Resid Ultrafining
VGO Ultrafining
RCD Isomax
Hydrobon Process
Scope of Application
Vacuum gas oil and lighter feedstocks
Whole crude, vacuum gas oil, and vacuum
tower bottoms
Vacuum tower residuum
Residual oil
High boiling virgin and cracked gas oil
Atmospheric tower residuum
Atmospheric tower residuum
Atmospheric tower residuum
Virgin or cracked heavy gas oils
Vacuum gas oil
Atmospheric tower residuum
Vacuum residuum
Vacuum gas oil
Atmospheric reduced crude, deasphalted
vacuum tower residuum
Light and heavy distillates
-------
(1) A Typical Hydrodesulfurization Process
A typical hydrodesulfurization process flow scheme, de-
rived from the Gulf HDS process, is illustrated in Figure XI-2.
This diagram reasonably represents most fixed bed residuum
desulfurization processes. Residuum desulfurization is the
process for the direct removal of sulfur for distillation column
bottom streams, commonly called residua. The atmospheric
reduced crude feed (that portion of the distillate remaining after
the most volatile fraction, gasoline, is removed at atmospheric
pressure) is supplied double desalted to limit the solids deposition
on the fixed bed catalyst. The reduced crude feed then passes
through a solids removal section which separates filterable
solids such as iron compounds. The oil feed is mixed with
makeup and recycle hydrogen, heated, and reacted over a
fixed bed of catalyst at elevated pressure and temperature where
H^S is evolved. Specific operating conditions vary depending on
the type of feedstock, the desired product, and the particular
process. For residuum feed, the pressure may range upward
from 71.2 kg/cm2 (1000 psig). Typical reaction temperatures
are 399 to 454 degrees C (750 to 850 degrees F). The reactor
effluent is cooled and recycle gas is separated. Sulfur is re-
moved from the recycle gas before it joins the reactor feed.
Separator liquids flow to fractionation or stripping.
When distillates are fed to desulfurization processes, the
solids removal step is unnecessary since these distillates are
already purified liquids. The remaining processing steps are
essentially the same as illustrated in Figure XI-2. However,
with lower boiling distillates the catalyst quantity needed in the
reactor is less and the pressure and temperature conditions are
less severe for a given level of sulfur removal. For distillates,
hydrotreating pressures may range upward from 29 kg/cm2
(400 psig) and the temperatures may be between 371 and 427 de-
grees C (700 and 800 degrees F).
An alternative system to the fixed bed reactor is used in
the H-Oil process which was developed by Cities Service R&D
and Hydrocarbon Research, Inc. This process uses an
ebullating bed of catalyst through which the oil and hydrogen
flow. The ebullating bed* process is competetive with fixed bed
Similar to a fluid bed but with liquid as the dispersing medium.
XI-11
-------
FIGURE XI-2
Typical Flow Scheme for a Fixed Bed
Hydrodesulfurization Process
Hydrogen Makeup
Atmospheric
Reduced
Crude
HDS Reactor
Section
Separation
Fractionation
Recycle Gas
Purification
H2S Rich Gas
Sour Gas*
-»»[ Distillates
Fuel Oil
* A gas fraction containing malodorous sulfur compounds.
XI-12
-------
systems primarily when residuums are fed. Advantages of the
ebullating bed reactor over a fixed bed reactor in such a system
are:
Solids contained in the feed pass through the
ebullating bed and do not cause plugging as occurs
with fixed beds
Catalyst can be added and removed from the ebullating
bed reactor while it is in operation, thus avoiding
the necessity for shutting down when catalyst
activity is too low
In general, smaller catalyst particles can be used
in ebullating beds. These show higher reactivity for
residuum desulfurization.
(2) Hydrogen Manufacture
A typical flow scheme for hydrogen manufacture from
natural gas by reforming is shown in Figure XI-3. In this figure,
natural gas feed enters the plant. Refinery gas could also be
used. Because catalysts used in reforming are rapidly deacti-
vated by sulfur, the natural gas is thoroughly cleaned. It may
be first scrubbed by caustic to eliminate acidic sulfur compounds
such as H?S. The natural gas may then flow through metal-
promoted activated carbon beds for removal of remaining traces
of acidic sulfur compounds and compounds such as COS and CS2-
The natural gas entering the reforming furnance should contain
less than about 0.1 ppm of sulfur. The reforming furnace feed
gas is preheated, mixed with steam, and passed through nickel
catalyst contained within furnace tubes. The reaction of steam
and methane takes place forming H2» CO, CO2» and a small
amount'of unreacted CH4 (a^O+bCH^^ cH2 +<*CO+PCO2+/CH4).
The tube skin temperature is 871 to 982 degrees C,(1600 to 1800 de-
grees F), and reforming pressure may range up to about 22.1 kg/
cm^ (300 psig). Product gases are quenched at the outlet of the
reforming furnace by injection of water and steam. The quenched
gas passes through an iron shift conversion catalyst to form ad-
ditional hydrogen (CO + H2O -» H2 + CO2). The shift conversion
reactor effluent is cooled, excess water, which is saturated
XI-13
-------
Steam >
Water and Steam
FIGURE XI-3
Typical Natural Gas Reforming Process
Natural Gas
Feed
Caustic Wash
I
Activated
Carbon Beds
Reforming Furnace
Catalyst Containing Furnace
Tubes
±
Shift Conversion
I
Separation
I
Acid-Gas Removal
Methanation
H2-Rich Gas to Compressors
Water
CO2-Rich Gas
XI-14
-------
with CO2, is separated, and the cooled gas stream passes
through acid-gas removal. Acid-gas removal may use the
monoethanol amine process, hot potassium carbonate pro-
cess, or one of the other CC>2 removal processes. Relatively
simple acid-gas removal processes may be employed because
only carbon dioxide need be removed. After CC>2 removal, the
main gas stream flows to methanation where the final traces of
CO and CC>2 contained in the gas (which may harm the HDS
catalyst) are converted to methane over a nickel catalyst
(aCO+ £CO2 + cH2 "* dCH4+ eH2O). The hydrogen-rich gas flows
to compressors for boosting to the required HDS reactor system
pressure. For the cases analyzed in this study, methane require-
ments to manufacture the process hydrogen are quantified below
in Table XI-2.
Table Xj-2
Hydrogen Plant Requirements and Production Rates
Methane required
(x 106ft3/day)
(m3/day)
Hydrogen manufactured
(x 106ft3/day)
(m3/dayj
Cases
1
1.65
46,725
3.30
93,450
2
12.00
339,850
24.00
679,700
3
16.00
453,125
32.00
906,250
3A
0
0
32.00
906.250
4
33.00
934,550
66.00
1,869,100
Alternative conventional hydrogen sources include hydrogen
from other refinery processes, naphtha reforming, or partial
oxidation of liquid hydrocarbons. Process hydrogen, which is
a by-product of another process, is the least expensive and the
most desirable if it is available. Important sources for this
hydrogen include excess gas from catalytic reforming which
contains hydrogen of purities from 65 to 95 percent depending
on the severity of the catalytic reforming operation. Hydrogen
from ethane cracking is recovered from the demethanizer. This
stream, which produces a gas stream that contains 85 to 90 percent
XI-15
-------
hydrogen, is suitable for use in HDS systems. Another potential
hydrogen source may be from acetylene plant by-product gas.
Many of these plants are based on partial oxidation of light hydro-
carbons. At the exit of the reaction zone, the gases are rapidly
cooled by direct quenching and the cooled gas is compressed
for purification. This gas contains 75 to 85 percent H2 plus
CO. After CO shifting, CO£ removal, and methanation of
residual CO and CO2* a high purity hydrogen can be produced.
Naphtha reforming is second to natural gas reforming as
the least expensive process for hydrogen generation. The :
naphtha reforming process is similar to natural gas reforming
as illustrated in Figure XI-3. The gasoline produced from crude
oil in the desulfurization Cases 1 to 4 could be used for generating
hydrogen via naphtha reforming.
Another conventional hydrogen manufacturing source is
partial oxidation using the Texaco or Shell processes. Liquid
hydrocarbon feed is injected into a reaction chamber in the
presence of steam with a minimum amount of oxygen for partial
oxidation. The resulting gases are quenched and purified.
Partial oxidation is the most expensive of conventional hydrogen
manufacturing systems discussed here but its versatility in
accepting a wide variety of feeds is a significant advantage.
Partial oxidation processes are discussed in greater detail in
another chapter of this report.
In the future, new processes for hydrogen manufacture
may be used. Several unconventional hydrogen processes are
presently being developed in conjunction with coal gasification
processes which require large quantities of inexpensive hydrogen.
One of these processes is gasification of petroleum coke with
steam and oxygen to produce a gas with a high content of hydrogen
and CO. The gas is shifted, acid-gas is removed, and the
residual CO and CO2 in the gas is methanated to produce the
product hydrogen. This process is similar to the coal and char
gasification process discussed in detail in other process studies.
* * * *
The five desulfurization schemes presented here are in order
of increasing severity of treatment. The discussion of pollution-
control processes is discussed in the next section:
XI-16
-------
(3) Case 1 (Figure XI-1 - Case 1)
Case 1 presents the simplest approach to distillate hydro-
desulfurization. Here, nc hydrodesulfurization of the residuum
is attempted. Gasoline represents 20.3 percent by volume of the
crude oil feed and boils below 150 degrees C (300 degrees F). It
is separated by distillation in an atmospheric topping unit and
routed elsewhere in the refinery. The remaining 79.7 percent
of oil enters the desulfurization section. Hydrogen for desulfur-
ization is assumed to be manufactured from natural gas, although
other hydrogen sources would be suitable. Naphtha, of 149 to
221 degrees C (300 to 430 degrees F) boiling range is produced
at a yield of 13.5 percent of the crude oil. The naphtha is pro-
cessed in a low pressure hydrotreater (a pressurized hydrode-
sulfurization vessel) which removes 95 percent of the feed sulfur
content. The gas oil from the atmospheric tower boils between
232 and 343 degrees C (450 and 650 degrees F). This gas oil,
representing 21.1 percent of the crude oil, is processed in sep-
arate hydrodesulfurization units because higher pressures and
larger amounts of catalyst are required to achieve the desired
sulfur reduction. Hydrodesulfurization of the gas oil reduces
feed sulfur content by 90 percent. In both the hydrotreater and
HDS units, sulfur is produced as hydrogen sulfide which is sep-
arated from the vented reactor gas. The total stabilized HDS
products are combined and blended with the 44 percent yield of
atmospheric residuum. This mix is called desulfurized oil. *
Vented gases from the hydrotreater and the HDS units
are combined and cleaned of ^S by monoethanol amine (MEA)
scrubbing. H2S-rich regenerated gas from the MEA treating
section flows to a Glaus plant for sulfur recovery. Tail-gas
from the Glaus is treated by one of several proprietary treating
processes such as the Beavon process. As seen from Figure XI-1,
Case 1, only 12 percent of the crude oil sulfur is recovered as
liquid sulfur from the Glaus unit.
Note that the hydrodesulfurized products may have greater
value as gasoline, etc., and, therefore, would not necessarily
be blended back into the residuum in most refineries. How-
ever, this blending is assumed in this study to permit process
comparisons.
XI-17
-------
In Case 1 a desulfurized oil sulfur content at 1.66 weight per-
cent is only slightly lower than the sulfur content of the crude oil
feed, 1.7 weight percent. However, the sulfur content of the oil
after removal of gasoline is 1.9 percent. The reason, of course,
is that about 85 percent of the total sulfur in the crude oil is con-
tained in the atmospheric residuum which is not directly desul-
furized. It is evident, however, that Case 1 alone cannot legiti-
mately be viewed as a desulfurization process.
(4) Case 2 (Figure XI-1 - Case 2)
Case 2 shows a process addition to Case 1 which will
partially remove sulfur from the residuum. In Case 2, topping
unit residuum is fed to a vacuum still which allows recovery of
31.8 percent of the original crude oil as vacuum gas oil distillate.
The vacuum gas oil is combined with atmospheric gas oil for
hydrodesulfurization; vacuum still bottoms represent 12.2 per-
cent of the crude oil and contain 42 percent of the crude oil sulfur
content. The desulfurizing products from hydrotreating and
hydrodesulfurization are combined and blended with vacuum
still residuum to form the desulfurized oil pool. The sulfur
content of this pool contains 0. 98 weight percent (slightly less
than 50 percent of the crude oil feed sulfur).
(5) Case 3 (Figure XI-1 - Case 3)
An additional processing complexity is added in Case 3 to
further reduce the desulfurized oil sulfur level. Solvent refining*
is added to separate sulfur containing extract from the vacuum
residuum. Because the extract contains less vanadium, nickel,
iron, carbon, sulfur, and nitrogen than the untreated residuum,
it can be hydrodesulfurized while the whole residuum cannot.
The solvent extract is processed in HDS Unit No. 2. The extract
is processed separately from the gas oil because it is a very
high molecular weight oil and requires more severe treating
A liquid-liquid separation process in which most of the asphaltenes
are separated from the light oil. The asphaltenes contain most of
the sulfur from the vacuum residual oil.
XI-18
-------
conditions than the atmospheric and vacuum gas oils. The HDS
No. 2 unit reduces the extract sulfur content by about 90 per-
cent. From 70 to 90 percent sulfur removal is considered good
for high boiling oils such ac vacuum resid extract. The desulfur-
ized products from the hydrotreater, from HDS No. 1 and from
HDS No. 2 are combined and blended with the solvent refining
unit raffinate.
The desulfurized oil pool for Case 3, at 0. 48 percent
sulfur, contains about 23 percent of the sulfur content of the
crude oil feed.
Additional process steam required for HDS processing
from the main refinery's boiler plant, and incremental process
fuel and electrical power requirements are reported in Table XI-3
for the more complicated cases analyzed in this study - Cases 3
and 4.
(6) Case 3A (Figure XI-1. Case 3A)
Case 3A is the same as Case 3 except that hydrogen is
generated by partial oxidation of the asphaltenes* (C + O% +
H2O—»H2 + CO 4- CO2 followed by a shift reaction) instead of
using natural gas steam reforming. Although greater capital
costs are encountered using partial oxidation and less oil product
is formed, the high sulfur raffinate not only makes hydrogen for
the HDS units but its consumption as partial oxidation feedstock
removes it from the desulfurized oil. Elimination of the raffinate
yields a desulfurized oil having 0.18 percent sulfur. Total
sulfur in the desulfurized oil represents approximately 10 percent
of the crude oil suflur content.
This approach toward total hydrodesulfurization, with no
residuum product, might be employed with crudes containing
high metals content. These crudes could not be processed by
direct desulfurization as discussed in Case 4.
The residual phase of a liquid-liquid solvent extraction process.
The other layer contains the solvent and extracted component.
XI-19
-------
Table XI-3
Steam Balance, Fuel and Power Requirements
N)
o
Case3
Process steam required
(x 103lb/day)
(x 1 03 kg/day)
Electrical power consumed (kWh/day)
Fuel required to generate process heat
(barrels/day)
Less plant production (FOE* barrels/day)
Net fuel purchased (barrels/day)
Case 4
Process steam required
(x 103lb/day)
(x 103 kg/day)
Electrical power consumed (kWh/day)
Fuel required to generate process heat
(barrels/day)
Less plant production (FOE* barrels/day)
Net fuel purchased (barrels/day)
Boiler
Plant
•-6.997
-3,174
19.871
1,623
-3,328
-1.510
9,434
771
Crude
Vacuum
Unit
4.179
1.896
1 2 1 ,000
1,800
2,902
1,316
84,158
1,152
Hydro-
treator
0
0
13,623
275
0
0
13,623
275
HDS
1,322
600
121,670
730
527
239
48,530
291
Solvent
Refining
and
HDS of
Extract
2,245
1,018
53,094
1,215
Resid
Desulfur-
ization
559
254
176,000
2,816
Sulfur
Recovery
-749
-340
2,130
174
-660
-299
1,890
155
Total
0
0
331,388
5,817
-307
5,510
0
0
333,635
5,460
-828
4,632
Fuel oil equivalent
-------
Note that use of the heaviest residuum from any of the
processes as partial oxidation unit feed will reduce the sulfur
content of the desulfurized oil, and eliminate the need for natural
gas for hydrogen manufacturing. The reason partial oxidation
has not been widely practiced in the past is because low cost
natural gas was available and the capital cost of natural gas
reforming was less. Partial oxidation, however, requires a
more complicated acid-gas removal process than the MEA
(monoethanolamine) systems employed in most refineries. This
point is discussed in more detail in another chapter of this report.
(7) Case 4 (Figure XI-1 - Case 4)
In Case 4, the crude oil is assumed to have low nickel and
vanadium content. Therefore, residuum can be directly desul-
furized economically. The nickel plus vanadium content of the
residuum feed to direct desulfurization must be 100 ppm or less
for economic operation. In Case 4 the atmospheric unit residuum
is fed directly to catalytic desulfurization for 90 percent sulfur
removal. This extent of sulfur removal represents the upper
limit of present day desulfurization capability for heavy feed-
stocks. The desulfurized oil for Case 4 contains between 9 and
10 percent of the total crude oil sulfur. Its sulfur content is
0.2 percent.
* >!< >!< # >;: *
As depicted in Figure XI-4, it is seen that as requirements demand
increased sulfur removal from the crude oil, a higher percentage of
the crude (more of the residuum) must be treated.
3. DISCUSSION OF POLLUTION CONTROL PROCESSES
The character and nature of the major waste streams of concern
in hydrosulfurization processes are discussed in this section. Control
treatment methods and alternatives for sulfur recovery are also
addressed.
(1) Sulfur Recovery
The sulfur which is produced from hydrodesulfurization of
oils is entirely hydrogen sulfide, H9S, and is recovered from reactor
XI-21
-------
FIGURE XI-4
Percent of Crude Oil Desulfurized for
Reduced Sulfur in Desulfurized Oil,
100,000 BPD Crude Oil, 1. 7 Percent Sulfur
8
S
a
LU
N
V)
UJ
a
OC
o
u.
o
d
(O
(O
o
CM
o
o
o
NI undins%'jj\A
XI-22
-------
vent gas. Some dissolved hydrogen sulfide is removed from
the liquid product by stripping and fractionating. In addition,
sulfur removal systems will be used within partial oxidation
plants such as is illustrated in Case 3A. It has been assumed that
the sulfur removal processes use monoethanolamine (MEA). The
MEA system i s a relatively simple acid-gas removal process that
has received wide application in petroleum refineries. It is
applicable in this case because the gas does not contain COS,
CS2* or HCN. Some of the other sulfur removal systems
which may be applicable include diethanolamine (DEA), Sulfinol,
hot potassium carbonate, Purisol, Rectisol, Stretford and
others. '-'
For most acid-gas removal processes, sulfur is con-
centrated as H2S which then must be converted to elemental
sulfur by processing through a Glaus unit. One-third of the
H2S feed to a Claus plant is burned in a boiler forming SO2t
The SO2 reacts with H2S over bauxite catalysts in one or more
reactor stages forming elemental sulfur. The sulfur is con-
densed and removed from the process as a liquid. The Claus
plant tail-gas contains not only unconverted H2S and SO2 but
also small amounts of COS, CS2» elemental sulfur vapor,
and elemental sulfur liquid mist. The sulfur content of the
tail-gas is reduced by a tail-gas treating process such as the
Beavon process, developed by Ralph M. Parsons Co. In the
Beavon process, the tail-gas is hydrogenated over a cobalt-
molybdenum catalyst (similar to HDS catalyst) to convert all
sulfur forms to hydrogen sulfide. After cooling, gas containing
hydrogen sulfide is treated by a Stretford system which converts
the H2S directly to elemental sulfur. Overall sulfur recovery
on the order of 99 percent can be achieved with this system.
The sulfur losses from this tail gas cleaning system in-
cludes:
Sulfates and thiosulfates which accumulate in the
recirculating liquor (requiring a small bleed
stream to wastewater treatment)
Sulfur dioxide which was not converted over the
hydrogenation catalyst
. Small quantities of hydrogen sulfide (perhaps 5 to
50 ppm).
XI-23
-------
In a refinery, this off-gas may be incinerated and could form part
of the oxidant (oxidizing agent) to support combustion for the re-
finery fuel.
(2) Gaseous Wastes
Although the Glaus tail-gas will be cleansed of most of its
sulfur, small amounts of sulfur (as K^S) will still be present in
the cleaned gas which can be incinerated (discussed in previous
paragraph).
Another waste gas stream is CC^-rich gas from the natural
gas reforming plant. The CC^ is formed during reforming and
is separated by an acid-gas removal unit within the hydrogen
manufacturing plant. Sulfur content of this gas should be nil
because of the extreme care used in cleaning the feed gas to the
plant.
Two other sources of waste gas are connected with the
operation of the hydrodesulfurization reactor section. The flows
of both gas streams are intermittent and are discussed below.
The activity of fresh HDS catalysts can be improved by
treating the catalyst with a sulfur-rich gas. Metals on the cata-
lyst are converted to sulfides. This activation procedure may
involve pressuring the system with hydrogen and the introduction
of H2S (or C$2 to form hydrogen sulfide) into the system for the
sulfiding reaction. During this step, some gas containing H^S
may be vented from the recycle gas system.
Another intermittent gas stream is the gas evolved when
deactivated HDS catalyst is regenerated by burning off carbon
deposits. One method of catalyst regeneration involves con-
trolling air injection into the HDS reactor system which is pres-
surized by inert gas. As carbon on the catalyst is burned off,
sulfur contained in the catalyst is evolved as SC^. The SC>2 is
scrubbed from the recirculating gas by injected caustic. How-
ever, some residual SC>2 may still remain in the gas which may
be vented to control the reactor pressure.
The fuel gas recovered during MEA scrubbing can be con-
sumed internally to meet process needs. The fuel-oil-equivalent
XI-24
-------
o
(FOE) heating value of this gas is 1.58 x 10 keal/barrel
(6. 3 x 10° Btu/barrel). Its composition can be found typically
to be:
Constituent
H
Cl
C2
C3
C4
C5
Total
%
36.7
47.0
12.0
3.8
0.3
0.2
100.0
The quantity of fuel gas produced for the cases analyzed in this
study are shown in Table XI-4 below:
Table XI-4
Fuel Gas Production, FOE barrels/day
Case
1
98
2
255
3&3A
307
4
828
(3) Liquid Wastes
One of the liquid streams is related to HDS catalyst re-
generation. This is spent caustic from scrubbing gases evolved
during catalyst regeneration. The spent caustic will contain
caustic, sodium sulfate, sulfite, possibly thiosulfates, and cer-
tainly carbonates and bicarbonates.
During normal operation of the HDS unit, its combined
gaseous and liquid effluent stream contains ammonium salts.
This ammonia is generated from hydrogeneration of the small
XI-25
-------
quantity of nitrogen in the initial feedstock (approximately
0. 1 kg nitrogen /bbl of crude) of the cases analyzed. Ammonium
salts reduce HDS catalyst activity and may plug downstream
piping. The ammonia can be removed by water-washing the
HDS effluent stream. The recovered water will contain dissolved
ammonia, H2S and CO2- This is the flow stream for which the
Chevron process was developed. Alternatively, this stream
could be stripped and the effluent gases directed to the sulfur
burner of a split-flow Glaus unit. The stripped water would
then contain approximately 15 ppm of E^S and 70 ppm of NH3;
this water is pure enough to be directed to the cooling water cir-
cuit. The expected water requirements for the four cases of
hydrosulfurization discussed are given in Table XI-5.
Table XI- 5
Expected Water Requirements
Process water required
(1/min)
(gal/min)
Case*
1
159
42
2
318
84
3
379
100
4
379
100
Case 3A would require additional water for the Texaco partial oxidation unit
as discussed in that process synopsis.
In the natural gas reforming plant, the shift conversion
effluent is cooled and water is condensed before the gases enter
the acid-gas removal unit. The shift converter effluent conden-
sate is saturated with CC^. There should be no sulfur contained
in this condensate except possibly on initial plant startup with a
new charge of catalyst. During initial startup, small amounts
of sulfur may be evolved from the catalysts and these might
appear in this stream.
Because natural gas reforming plants operate at modest
pressures, the product gas must be compressed to the high pres-
sure of the HDS systems. Several stages of compression must
be used. Condensate collected at the compressor interstage
coolers will not only contain some dissolved gases, but also
small amounts of lubricating oil from the compressors.
XI-26
-------
(4) Solid Wastes
The solids produced from the desulfurization unit are the
spent catalysts from the different sections of the plant. HDS
catalyst must be replaced when poisoned by vanadium,iron and
nickel compounds. When contaminated with carbon, the catalyst
can be regenerated by burning. The spent catalyst will contain
cobalt, molybdenum, nickel, alumina, and the contaminants
vanadium, nickel, and iron. Before removing the catalyst from
the reactors, it should be oxidized under controlled conditions
to remove the pyrophoric (capable of igniting spontaneously) iron
sulfide and carbonaceous deposits.
Several different solids products will be available for dis-
posal from the natural gas reforming. One solid is the spent
metal-promoted active carbon, which can be regenerated several
times by steam and air. During regeneration, the adsorbed sulfur
compounds convert to elemental sulfur which is then readsorbed
by the carbon. When the sulfur content of the carbon reaches
70 percent or more, its activity declines and it must be dis-
carded. The catalyst should be regenerated before disposal to
convert all of the sulfur to the elemental form.
The natural gas reforming plant contains at least three
kinds of catalysts. The reforming catalyst, which is contained
in the reforming furnace tubes, is nickel on a refractory support.
The shift catalyst is iron oxide and the methanation catalyst is
nickel on alumina. The nickel catalyst is very sensitive to sulfur
and can be quickly deactivated when small amounts of sulfur are
present in the processed gas. More frequently, reforming cata-
lyst is deactivated by loss of the necessary ratio of steam to
hydrocarbon feed. This causes excessive carbon deposition on
the catalyst. The catalyst must then be removed from the furnace
tubes and replaced with fresh catalyst. Over a period of time,
iron shift conversion catalyst gradually decreases in activity,
probably as a result of changes in its crystalline structure. In-
advertent injection of liquid water into a bed of hot shift catalyst
will cause catalyst disintegration, thereby increasing the pres-
sure drop across the catalyst bed, and requiring replacement of
the catalyst. Methanation catalyst is rarely deactivated by sulfur
because of its location in the process. However, the methanation
catalyst is subject to variations in carbon oxide content of its
feed. When the total carbon oxides exceed about 1. 5 percent in
XI-27
-------
the methanator feed, runaway temperature rises occur, causing
catalyst deactivation. The reforming, shift, and methanation
catalysts should be oxidized within the plant under controlled
conditions to return them to a stable form before removal.
Otherwise, the catalysts are pyrophoric. Spent reforming and
methanation catalysts from large plants will probably be re-
turned to the catalyst supplier for recovery of nickel. The iron
catalyst will probably be discarded.
Bauxite is used in the Glaus system as a catalyst. The
bauxite may require replacement as a result of plant opera-
tional problems, the most frequent of which are:
Unexpectedly high-hydrocarbon content in the feed
gas and a deficiency in the amount of combustion air.
This causes formation of soot which is deposited in
the bauxite bed causing plugging.
Excessive amounts of steam introduced to the bauxite
bed which tend to soften the catalyst. This causes
increased pressure drop and requires catalyst re-
placement.
Another problem which occasionally occurs is con-
densation of sulfur within the catalyst bed because of
poor catalyst bed temperature control. The sulfur
liquid plugs the bed creating an increased pressure
drop and, if the bed is badly plugged, requires cata-
lyst replacement.
Spent bauxite, therefore, may contain soot and sometimes
may also be contaminated with elemental sulfur. Bauxite and its
contaminants are stable so they can be disposed of as landfill
without further treatment.
The estimated amounts of catalyst exhausted during plant
turnarounds (perhaps an annual occurrence) are quantified in
Table XI-6.
XI-28
-------
Table XI-6
Spent Catalyst, Ib/day (kg/day)
Unit Process
Hydrotreating
HDSofgasoi!
HDS of vacuum residuum
Direct desulfurization of residuum
Case
1
185 (84)
243(110)
-
-
428(194)
2
185 (84)
487(221)
-
-
672(305)
3
185 (84)
485(220)
952(432)
-
1622(736)
4
185 (84)
243 (110)
-
8849(4014)
9277(4208)
4.
COSTS OF POLLUTION CONTROL*
The costs for pollution control by hydrodesulfurization of crude
oil fractions are presented in this section (the format is derived in
Chapter III of this report). They are 1973 costs for East Coast U. S.
refineries. Cost information was obtained from process licensors,
literature sources, and from discussions with technical advisors who
are experienced in refinery processing. The costs presented are
based on sulfur removal only. The cost for treatment of other wastes,
generated by the desulfurization section (discussed in Section 3), cause
incremental effects in existing waste treatment facilities and therefore
are not considered.
The incremental** capital costs for pollution control are reported
in Table XI-7 for Cases 1, 2, 3, and 4. Capital costs are separated
into costs directly related to hydrodesulfurization of oil and costs
directly related to sulfur recovery from fuel gases generated by the
hydrodesulfurization processes. These costs are incremental since
it was assumed that the sulfur control processes are added to an exist-
ing petroleum refinery.
The costs presented here are preliminary estimates; the actual
costs of emission control in this process will not be known until
a full-scale facility has been constructed and operated.
Incremental costs in this report refer to that portion of the total
plant cost that is assignable to pollution control facilities.
XI-29
-------
Table XI-7
Incremental Capital Investment3- ($ Million)
Item
Incremental plant investment
Vacuum distillation
Solvent refining
Hydrotreatcr
Light distillate desulfurizer
Heavy distillate desulfurizer
Deasphalted oil desulfurizer
Reduced crude desulfurizer
Hydrogen manufacture
Boiler facilities
MEA scrubbing
Claus unit
Claus tail gas unit
Subtotal plant investment
Project contingencies4"
Incremental plant investment (I)
Startup costd (S)
Interest during construction6 (IDC)
Incremental working capital (W)
Incremental capital investment (C)
Case 1
Desulfurized
Oil
_
6.0
10.0
-
6.4
4.0
-
-
26.4
4.0
30.4
0.9
7.2
0.8
39.3
Clean Fuel
Gas
._
-
-
—
1.2
0.4
0.4
2.0
0.3
2.3
0.1
0.5
0.1
3.0
Case 2
Desulfurized
Oil
7.2
6.0
10.0
11.4
12.0
11.0
-
-
57.6
8.6
66.2
2.3
15.7
1.8
86.0
Clean Fuel
Gas
_
-
-
-
3.3
1.1
I.I
5.5
0.8
6.3
0.2
1.5
0.2
8.2
Case 3
Desulfurized
Oil
7.2
10.6
6.0
10.0
11.4
17.4
13.8
12.4
-
-
88.8
13.3
102.1
3.6
24.2
2.7
132.6
Clean Fuel
Gas
_
-
-
—
4.5
1.5
1.5
7.5
1.1
8.6
0.2
2.0
0.2
11.0
Case 4
Desulfurized
Oil
_
6.0
10.0
45.4
21.4
7.2
—
-
90.0
13.5
103.5
3.2
24.5
2.7
133.9
Clean Fuel
Gas
_
-
-
—
4.8
1.6
1.6
8.0
1.2
9.2
0.2
2.2
0.2
11.8
CO
o
Notes:
Incremental plant investment, return on investment during
construction, and working capital are treated as capital costs
in the year 0 (the year ending with completion of startup
operation).
Installed costs, including engineering design costs, contractors'
profit, and overhead.
Includes costs for unexpected site preparation and hardware
requirements at 15% of plant investment.
d At 20% of incremental gross operating cost.
e Computed as the discount rate x incremental plant investment
for 1.875 years' average construction period.
I(l-H)n = I(1.12)'-875 = 1.236761 or 0.2367611 additional investment.
f Sum of materials and supplies at .9% of incremental plant investment
and net receivables at 1/24 of annual incremental revenue.
-------
Since hydrodesulfurization process installations will be constructed
in the private sector, cost calculations have been made using only the
DCF financial accounting method.
The incremental capital investment rapidly increases as the oil
is further tested to reduce its sulfur content (see Table XI-8 and Fig-
ure XI-5). The mildest hydrodesulfurization scheme analyzed (Case 1)
costs $42.5 million; the severest (Case 4) requires $145.7 million.
Tables XI-8, -9, and -10 present the incremental capital and op-
erating costs for each of the cases. The operating costs are also sepa-
rated into costs directly related to hydrodesulfurization and costs di-
rectly related to fuel gas cleanup.
The incremental annual operating costs increase substantially
as the desulfurized oil quality is improved. For a 100,000 barrel/day
1. 7 percent sulfur crude oil plant to decrease the sulfur content of the
feed by about 0. 6 percent (by weight) of the oil has the effect of
doubling the annual operating cost of desulfurization (see Figure XI-6).
The derivation of the formulae for calculating the incremental
cost of pollution controls is presented in Chapter III. For the DCF
method, the required incremental annual cost, X, for the assumed
rate of return is:
X = N + 0.238161 + 0. 1275S + 0. 23077W
where: N = incremental net operating cost
I = incremental plant investment
S = startup costs
and W - incremental working capital.
The annual crude oil feedstock supplied (H) is
(100,000bbl/day)x 0.9 capacity factorx 365 days/yr= 32.85xl06bbl/yr.
Therefore, the incremental cost of desulfurizing crude oil, when
expressed as $/bbl, is
X X
H 32.85x 106 bbl/yr
XI-31
-------
K)
O
Q
UJ
N
E
D
O
z
tc
D
o>
1.8
1.6
1.4
1.2
1.0
0.8
0.6
0.4
0.2
0.0
. FOR CLEAN FUEL GAS
FORDESULFURIZEDOIL
20
40
60
80
100
120
140
Sn
P-8
o 5-'
o P
o ^
w 51
*"lj ^
O 03
c-t- "T]
O B H-i
21 §
a 5- S
fC ?U
I
CJ1
INCREMENTAL CAPITAL INVESTMENT, $10°
o
0)
o
e-l-
(-"•
o
-------
Table XI-8
Incremental Annual Operating Cost ($)
.
W-'.
Item
Labor
Direct operating labor
Maintenance labor(1.5% of 1)
Supervision ( 1 57c of direct operating
and maintenance labor)
Incremental labor cost
Administration and general overhead
(6V7i of incremental labor)
Other direct costs*
Supplies
Operating (30% of direct operating labor)
Maintenance (\.57c of I)
Incremental cost of supplies
Local taxes and insurance (2.7% of I)
Incremental gross operating cost
By-product credits
Incremental net operating cost (N)
Case I
Desulfurized Clean Fuel
Oil Gas
309.000 103.100
456.000 34.500
114.800 20.600
879.800 158.200
527.900 94.900
1.877.700 1.300
92.700 30.900
456.000 34.500
548,700 65.400
820.800 62.100
4.654,900 381.900
4.654,900 277,800
Case 2
Desulfurized Clean Fuel
Oil Gas
538.400 182.300
993.000 94.500
229.700 41.500
1.761.100 318.300
1.056.700 191.000
5.549.700 4.600
161.500 54.700
993.000 94.500
1.154,500 149.200
1.787.400 170,100
11.309.400 833.200
1 1 .309.400 398.900
Case 3
Desulfurized Clean Fuel
OU Gas
847,300 236,000
1.531,500 129.000
356.800 54.800
2.735.«00 419,800
1.641.400 251.900
8.933.600 9.200
254.200 70,800
1.531.500 129,000
1.785.700 199,800
2.756.700 232,200
17,853.000 1.112.900
17.853.000 475.300
Case 4
Desulfurized Clean Fuel
Oil Gas
540,100 257,400
1,552,500 138,000
313,900 59,300
2,406.500 454,700
1,443,900 272,800
7,485,800 7,900
162,000 77,200
1,552,500 138,000
1,714,500 215^200
2,794,500 248,400
15,845,200 1.199,000
1 5,845,200 448,000
* Includes catalysts, chemicals and utilities
-------
Table XI-9
Summary of HDS Capital Costs
Case
1
2
3
4
Desulfurized Oil
(Wt % Sulfur)
1.66
0.98
0.48
0.20
Desulfurized Oil
Incremental
Capital Costs
($ Million)
39.3
86.0
132.6
133.9
Desulfurized Fuel Gas
Incremental
Capital Costs
($ Million)
3.0
8.2
11.0
11.8
Table XI-10
Summary of HDS Operating Costs
Case
1
2
3
4
Desulfurized Oil
(Wt % Sulfur)
1.66
0.98
0.48
0.20
Desulfurized Oil
Incremental Annual
Operating Costs
\X$/bbl of Crude Oil)
0.37
0.85
1.32
1.26
Desulfurized Fuel Gas
Incremental Annual
Operating Costs
($/bbl of Crude Oil)
0.03
0.06
0.08
0.08
XI-34
-------
8
I
OJ
en
o,
o
ui
N
DC
UJ
Q
D
LU
D
1.8
1.6
1.4
1.2
1.0
0.8
0.6
0.4
0.2
0.0
FOR CLEAN FUEL GAS
FOR DESULFURIZEDOIL
I
I
.20
.40
.60 .80
INCREMENTAL OPERATING COST, $/BBL
1.00
1.20
1.40
O
o
00
c-K
O •->
o
cj
f\ "> I
8 o »
C3
^^ ^—4
g1 §
>-• *
i~p o
5 o
""* o
W
hd
a
-------
Solutions for the above relationships are shown in Table XI-11. As
practiced in the industry, crude oil feedrate quantities were used to
express incremental costs. In addition, to facilitate process com-
parisons, cost per Btu were computed using the conversion factor:
1 bbl crude = 5. 875 x 106 Btu
(19, 480 Btu/lb) (7. 18 Ib/gal) (42 gal/bbl) = 5. 875 x 106 Btu/bbl.
Table XI-11
Incremental Cost of Desulfurizing Crude Oil
Incremental annual cost (X). $/yr
Incremental cost of desulfurizationf— ), $/bbl
\x/
Incremental cost of desulfurization:
$/bbl * 5.875 x 106 Btu/bbl, $/Btu
Case 1
Desulfurized
Oil
12.194.400
0.37
Clean Fuel
Gas
861.400
0.03
0.07
Case 2
Desulfurized
Oil
27.784,400
0.85
Clean Fuel
Gas
1,971,000
0.06
0.15
Incremental annual cost (X), $/yr
Incremental cost of desulfurization ( — j, $/bbl
\ A./
Incremental cost of desulfurization:
S/bbl -^ 5.875 x 106 Btu/bbl. $/Btu
Case 3
Desulfurized
Oil
43.251,400
1.32
Clean Fuel
Gas
2,595,100
0.08
0.24
Case 4
Desulfurized
Oil
41,526,100
1.26
Clean Fuel
Gas
2,710,700
0.08
0.23
Before directly comparing these costs with those derived for the coal
gasification processes investigated as part of this study, it should be
cautioned that the crude oil feedrate of 100, 000 bbl/day, dictated by
present-day oil refinery economies, constitutes 150 percent of the
energy input used for the coal gasification processes. It should also
be emphasized that hydrosulfurization of crude oil does not produce
an end product fuel oil. It simply supplies a low sulfur pool of oil
which must be subsequently refined to yield any of numerous fuels
derived from distillation and blending.
XI-36
-------
XII. THE GAS COMBUSTION RETORT PROCESS
Oil shale is an impermeable, sedimentary rock containing solid,
organic material called kerogen. When heated, kerogen can
be transformed into liquid and gaseous hydrocarbons, which can be
further processed to make fuel oil. Oil shale deposits in the Rocky
Mountain area of the United States (Colorado, Utah, and Wyoming)
contain about 600 billion barrels of this type of oil. Interest in the
commercial development of this extensive potential source of energy
has fluctuated widely over the last 100 years. Some oil from shale
was produced prior to the 1859 discovery of natural petroleum, but
no serious attention was focused on oil shale until just prior to 1920.
At that time there was concern that domestic petroleum resources
might not be adequate to meet anticipated demand. New petroleum
discoveries, however, resulted in declining interest until 1944 when
a Synthetic Liquid Fuels program was established due to the demand
for liquid fuels during World War II. The present energy crisis again
has renewed interest in producing oil from shale. Many private
companies and the Bureau of Mines are working to improve existing
oil shale-to-oil conversion processes and to develop new ones.
Two major options are being considered:
. Strip or deep-mining followed by surface processing of
the oil shale
. In situ (in-place) oil;recovery, that is, underground
processing of the shale.
Of these options, only the first approach has been developed to the
point where scaleup to commercial production is expected in this
decade. The in situ approach is still in the experimental stage and
will require significant development work before being commercially
viable. In this chapter, only surface processing will be considered.
Oil is recovered from oil shale by heating it to between
425 degrees and 540 degrees C (800 degrees to 1000 degrees F) in
a retort (a heating vessel which decomposes the shale) either directly
or indirectly. The following three retorting processes are considered
XII-1
-------
to be candidates for early commercialization from the several
processes under development. The processes are the:
Union Oil Retort process
TOSCO II Retort process
Gas Combustion Retort process.
Oil shale retorts are essentially heat exchangers for trans-
ferring heat from a heating medium to the oil shale. For the Union
Oil Retort and Gas Combustion Retort processes, heat is trans-
ferred to the shale directly from the combustion occurring in the
retort by burning product gases and residual carbon in the retorted
shale. In the TOSCO II Retort process, heat is transferred to the
shale indirectly by the introduction of hot solids (ceramic balls) into
the retorting bed. Since information was relatively complete for
the Bureau of Mines' Gas Combustion Retort process, it was used as
the basis for this synopsis.
The Bureau of Mines designed, built, and operated three pilot
plants (of 5, 23, and 136m ton/day capacity) at Anvil Points near
Rifle, Colorado, until operations were discontinued in 1955.
1. PROCESS DISPLAY
Figure XII-1 contains the process flow diagram for a facility
with a 50, 000 bbl/day oil production capacity. This is equivalent
to about 70 x 109 kcal/day (277 x 109 Btu/day). The associated
energy balances, flow rates, and stream compositions for this
facility are also contained.
The data available on the Gas Combustion Retort process were
collected before the need for environmental considerations was
emphasized. As a result, the data available on potential pollutants
are sparse. In this synopsis, therefore, an attempt has been made
to describe a process based on a hypothetical average shale. These
data may have to be modified for a particular plant installation.
(1) Basis for Analysis
The process flow sheet data was based on information
developed by the Bureau of Mines from operation of its Gas
XII-2
-------
FIGURE XII-1
The Gas Combustion
Retort Process
65,960 m ton/day
50.000 btt/dav (7.95 X 10° I/day
X
LEGEND
LFE
MTONS
L/OAY
M
S
SUS
LIQUID FUEL EQUIVALENT
HEATING VALUE OF 6.3 X 1
(9,987 KCAL'L)
METRIC TONS
LITERS PER DAY
METERS
SULFUR
WITH
6 BTU/BBL
SAYBOLT UNIVERSAL SECONDS
STREAM NO.
STREAM NAME
COMPONENT
HEATING VALUE
®
OIL SHALE
WT.*
ORGANIC MATTER 13,8
C 80.5
H 10.3
N 2.4
S 1.0
0 5.8
TOTAL 100.00
MINERAL MATTER 86.2
CARBONATES 48.0
FELDSPARS 21.0
QUARTZ 13.0
CLAYS 13.0
ANALCITE 4.0
PYRITE 1.0
TOTAL 100.00
2520BTU/LB (1400 KCAL/KG)
®
SPENT SHALE
WT.%
ORGANIC MATTER UNK
Si O2 42.3
Fe2O3 4.5
AI2°3 )3°
C«O 23.1
MgO 9.9
SO3 1.8
KjO 2.3
TOTAL 100.0
786 BTU/LB (159 KCAL/KGI
©
RETORT GAS
MOLE*
N2 62.1
CO2 24.5
HjS 0.1
H2 5.7
HYDROCARBONS 5.3
TOTAL 100.0
FLOWRATE24.2X 106FT3'HR.
I685.3X 103M3/HR.I
100 BTU/FT3 (890 KCAL/M3!
©
RAW SHALE OIL
WT.%
S 0.74
N 2.18
GRAVITY °API, 19.7
POURPT, 27°C
VISCOSITY 9 38°C. 2256 SUS
'
©
UPGRADED SHALE
WT.%
S 0005
N 0035
GRAVITY °API, 46.2
POURPT.OO°C
VISCOSITY e 38°C, 40 SUS
' '° ° ° L/KG'
ENERGY BALANCE
CARRIER
OIL SHALE
TOTAL INPUT
SHALE OIL
SPENT SHALE
COKE
AMMONIA
SULFUR
LOSSES
TOTAL OUTPUT
X106BTU/HR. (X106KCAL MR.l
15477.0 43900)
15477.0(3900)
11559.0(2912.51
1405.2 13541
1004.1 (2S3)
111.2(28)
23.4 (5,9)
1374^ (346.3)
15477.0 139001
0Y THE PROCESS DEVELOPER
-------
Combustion Retort pilot plants at Rifle, Colorado; from the
references cited at the end of this process synopsis; and from
private communication with the process developer.
The analysis shown is for a 0. 6 percent sulfur content oil
shale from the Mahogany Zone of Colorado and Utah. This shale
contains 10 to 20 percent organic matter, the balance being
mineral matter. The composition of this raw shale feed is
given in Table XII-1.
Table XII-1
Oil Shale Component Analysis
Major Components
Organic matter
Subtotal
Mineral matter
Subtotal
Total
Wt.%
13.8
82.6
100.00
Subcomponents
Carbon
Hydrogen
Nitrogen
Sulfur
Oxygen
Carbonates
Feldspars
Quartz
Clays
Analcite
Pyrite
Weight %
80.5
10.3
2.4
1.0
5.8
100.00
48.0
21.0
13.0
13.0
4.0
1.0
100.00
XII-4
-------
The heating value of the organic matter is 1400 kcal/kg
(2520 Btu/lb). It is assumed that 125 liters of oil per m ton
of oil shale (30 gal/short ton) are produced, along with 1. 94
of retort gas per liter of oil produced (10, 900 ft^/bbl) with a
heating value of 890 kcal/m3 (100 Btu/ft3). This gas is used
as plant fuel. The oil-shale to shale-oil process thermal
efficiency is determined to be 74. 7 percent.
(2) Layout and Symbols
The general direction of process flow is from the shale
supply on the left to the snycrude (synthetic crude oil) on the
extreme right of the flow diagram. Residuals and by-products
are shown along the bottom. Inverted trapezoids denote residual
storage. The bold line indicates the flow of the primary oil
producing process.
The circled figures refer to the stream compositions
which also appear as tabulations in Figure XII-1. The overall
energy balance is also shown there and the energy balance
calculations are given in Table XII-2.
Rhombic-shaped units represent intermediate products
or sources for which the distribution is not shown. Nonintegral
cleanup processes are indicated by sloping rectangles. These
treatment and recovery processes are discussed in the remain-
ing sections of this chapter. For a more thorough explanation
of each licensed cleanup process mentioned, refer to the dis-
cussion given in Chapter III as well as to readily available
process literature and the bibliography included in this report.
2. PROCESS DESCRIPTION
The unit processes which comprise the Gas Combustion Retort
scheme are described in this section. By -product manufacture and
pollutant generation are discussed where appropriate. The dis-
cussion of pollution control processes is covered in the next section,
XII-5
-------
Table XII-2
Energy Balance Calculations for the
Gas Combustion Retort Process
I
ON
Carrier
OU shale
Total Energy Input
Product shale oil
Spent shale
Coke
Ammonia
Sulfur
Losses*
Total Energy Output
Calculations
73,700 short tons/day x 2000 Ib/short tons x 2520 Btu/lb+24 hrs/day
50,000 bbl/day-i-24 hrs/day x 42 gal/bbl x 6.63 Ib/gal x 19,925 Btu/lb
58,960 short tons/day x 2000 Ib/short ton-r-24 hrs/day x 286 Btu/lb
855 short tons/day x 2000 Ib/short ton-i-24 hrs/day x 14093 Btu/lb
138 short tons/day x 2000 Ib/short ton-i-24 hrs/day x 9668 Btu/lb
3.52 short tons/hr x 2000 Ib/short ton x 3983 Btu/lb
(by difference)
(x 106 Btu/hr)
15,477.0
15,477.0
11,559.0
1,405.2
1,004.1
111.2
28.0
1,369.5
15,477.0
(x 106 kcal/hr)
3900
3900
2912.8
354.0
253.0
28.0
7.1
345.1
3900
Includes sensible heat of product, by-products, and waste streams as well as air cooling and other incremental losses.
-------
(1) Shale Preparation
The shale is conveyed directly to the receiving hoppers
at the crushing plant by truck or conveyor (Stream (A) ). It is
crushed to minus 26. 7 cm (10. 5 inch) size in the primary
crushers. Secondary crushers reduce the size to minus 7. 62 cm
(3 inch). The crushed shale passes through screens to remove
minus 0.5 cm (3/16 inch) particles for briquetting before going
to the retorts. Particles between 2.54 and 7.62 cm are con-
veyed to surge bins before retorting. According to Bureau of
Mines report^), 1.3 percent of the shale handled during
crushing and screening is lost as dust. Half of this loss is
assumed to occur during crushing and transporting and the
balance in screening (sizing).
(2) Retorting
The following detailed description of the operation of a
retort is based on that reported by the USBM (ref. 2). Some
of the operations described occur within the retort and as such
are not depicted. The crushed and sized or briquetted shale
from the surge bins is fed to the retort feed hoppers atop the
retorts by belt conveyors equipped with automatic trippers to
feed the individual retorts. For a 7,948,500 I/day (50,000 bbl/
day) plant, six 17 m diameter retorts having a shale bed depth
of 5. 5 m are required. The fresh feed at the top of each unit
is preheated by the off-gases from the retort combustion zone.
In the combustors, located near the midpoint of the shale bed,
recycled low-Btu gas is burned with air to heat the shale to the
maximum temperature of about 927 degrees C to 980 degrees C.
Approximately 82 percent of the recycle gas is fed to the bottom
of the retort and is utilized to cool the spent shale to about 93 de-
grees C prior to discharge. The remainder of the recycled gas
with the combustion air is fed directly to the combustor.
Within the retorting plant, the evolved gas, containing
entrained crude shale oil, flows through impingment type
separators, centrifugal separators and electrostatic precipi-
tators to separate the gas and oil. The retort gas has a heating
value of about 890 kcal/m3 (iQO Btu/ft3 as shown by Stream (T)
of Figure XII-1. It is compressed for recycle to the retort
and for use as fuel elsewhere within the process. By petroleum
standards, the raw shale oil (as described in Stream (1T))is a
low gravity, high nitrogen, moderate sulfur crude oil. It
XII-7
-------
can be further refined by standard crude refinery procedures.
The refining scheme used for this analysis is contained within
the "Shale Oil Upgrading Sections, " as shown in Figure XII-1.
(3) Shale Oil Upgrading
Upgrading of the retorted raw shale oil (Stream (¥) ) is
accomplished through distilling, delayed coking and hydro-
genating unit processes which are standard petroleum refining
operations. The raw shale oil is first heated and sent to a
distillation column where it is separated into approximately
50 percent light and 50 percent heavy oil fractions. The light
fraction (overhead vapors) is cooled to yield a distillate product,
and sent to the hydrogenation unit. The heavy fraction (bottoms)
from the distillation column is fed through a heater and sent to
the delayed coking unit at a temperature of 504 degrees C.
Coking is a process in which the raw feed is subjected
to destructive distillation to yield a solid residue called coke
as well as organic liquid and gas products. In delayed coking
by increasing liquid phase residence time and not quenching,
(allowing higher coking temperatures), the amount of oil (liquid
distillate) evolved is increased. The liquid distillate is cooled,
depropanized* and sent to the hydrogenation unit. Coke from
delayed coking (776 m tons/day) is stored for sale.
The combined distillate from distillation and delayed
coking is hydrogenated (cracked to simpler hydrocarbons by
heat in the presence of hydrogen). Hydrogenation takes place
at 446 degrees C (835 degrees F), a pressure of 106.5 kg/cm^
(1500 psig) and requires 0. 3 m3 of hydrogen per liter of distillate
feed (1700 ft3/bbl). Though it is not specifically stated by the
Bureau of Mines, the low sulfur and nitrogen content of the
upgraded oil indicates that catalytic hydrogenation was employed.
The properties of this upgraded oil are shown as Stream (In in
Figure XII-1.
The gas streams from the hydrogenation and delayed
coking units contain recoverable sulfur and nitrogen. In the
* The gaseous portion (that is, propane, CgHg and lighter
fractions) is removed.
XII-8
-------
Gas Combustion Retort process, as developed, these streams
are water-washed twice in the gas treating unit. The first
wash employs an ammonia solution which acts as an acid gas
treatment, removing H2S. The second wash uses water to
remove the remaining ammonia. The sulfur and nitrogen is
stripped and extracted as an H2S-NH3-water stream. This
combined stream is separated in a recovery unit where, by
heating to 77 degrees C (170 degrees F), the H2S is driven off.
Traces of ammonia are removed from this H2S stream by
sulfuric acid scrubbing. A Glaus unit with a Beavon tail gas
cleanup system is used in this analysis to recover sulfur from
the H2S. Note that this tail gas treatment would be incorporated
into the Wellman-Lord system used on the power plant effluent.
The ammonia solution is then pressurized to 17. 2 kg/cm^
(230 psig) and heated to 166 degrees C (330 degrees F) to
liberate the anhydrous ammonia which is cooled, condensed
and stored as a liquid for sale. The water is recirculated to
the gas treating unit.
(4) Hydrogen Generation
About 89 percent of the washed gas from the gas treating
unit is used to produce the hydrogen required for hydrogenation.
Hydrogen is produced by steam reforming the gas at 4. 5 kg/cm2
(50 psig) and 760 degrees to 816 degrees C (1400 degrees to
1500 degrees F) over a nickel catalyst, followed by catalytic
water gas shifting (CO + H2O-»-H2 + CO2> at 427 degrees C
(800 degrees F), purification and compression. The remaining
11 percent of the gas, along with the retort gas, is used as a
plant fuel.
(5) Steam and Power Generation
The steam and electric power requirements for the
process are met by an on-site steam and power plant fired by
the retort gas and the fuel gas from oil upgrading sections (the
gas treating unit). The resultant stack gases contain significant
SO2, about 2. 9 kg/106 kcal (1. 6 Ib SO2/106 Btu) of fuel burned.
The proposed sulfur limit in the effluent from fuel gas com-
bustion is equivalent to 23 gms/100 m3, (10 grains sulfur/100 ft3)
of fuel gas, requiring 83 percent stack gas cleanup. A Wellman-
Lord system was arbitrarily applied for stack gas cleanup
XII-9
-------
in this analysis; this system should be capable of 90 percent
removal.
(6) Energy Balance
The overall energy balance for the process is presented
in Figure XII-1 and derived in Table XII-2. From this energy
balance, the oil shale-to-shale oil process efficiency is found
to be:
2913 x 106 kcal/hr^ 3000 x 106 kcal/hr = 74.7%
(7) Sulfur Balance
The sulfur balance for the 0. 6 percent sulfur content oil
shale feedstock is quantified in Table III-3 and discussed in
Section 3 of this chapter. Of the 16, 683. 4 kg/hr of sulfur
Table XII-3
Sulfur Balance for the Gas
Combustion Retort Process
Carrier
Short tons/hr
kg/hr
Raw oil shale
Organic sulfur
73,700 short tons/day + 24 hr/day x . 138 x .01
Pyritic sulfur 64.132S-,
73,700 short tons/day 7 24 hr/day x .862 x.01 x —
119.982FS2
Total input
4.24
14.15
18.39
3,846.5
12,836.9
16,683.4
Elemental sulfur
84.4 short tons/day i 24 hr/day
Spent shale
Pyritic sulfur (as above)
Organic sulfur
Product shale oil
Sulfur vented to atmosphere
From sulfur recovery
From stack gas cleanup
Total output
3.52
14.15
0.50
0.01
0.04
0.17
18.39
3,193.3
12,836.9
453.6
9.1
36.3
154.2
16,683.4
XII-10
-------
carried in the raw feed, about 77 percent appears in the
mineral matter as pyrite (FeS2), while the remaining 23 per-
cent is organic sulfur. Only trace amounts of sulfur are
present as sulfates.
All of the pyritic sulfur and about 12 percent of the organic
sulfur reports to the spent shale which serves as minefill. Of
the remaining 3392. 9 kg/hr of sulfur, about 94 percent is
recovered as elemental sulfur for sale, about 6 percent is vented
to the atmosphere from the Glaus and Beavon sulfur recovery
and Wellman-Lord stack gas cleanup plants, and less than 0.03
percent remains with the product shale oil.
3.
DISCUSSION OF POLLUTION CONTROL PROCESSES
The major waste streams of concern in the Gas Combustion
Retort process for producing oil from shale, and their proposed
control and treatment methods are summarized in Table XII-4.
(1) Shale Dust
Most of the shale dust is produced during the storing,
crushing, and screening operations. The conventional
Table XII-4
Nature and Treatment of Major
Waste Streams
Final Wastes
Sources
Treatment
Shale dust and particulates
Spent shale
Wastewater
Hydrogen sulfide
(intermediate stream)
Sulfur dioxide
Shale storage, crushing,
briquetting, conveying
Retorting plant
Mining, retorting, oil
upgrading
Distillation, delayed
coking, hydrogenation
Sulfur recovery, steam
and power plant
Cyclone separators,
bag filters, enclosures
Surface and mine
disposal
Biological and modified
Chevron process, return
to river and mine
Claus and Beavon
processes
Wellman-Lord and
Beavon processes
XII-11
-------
technology of using cyclone separators, bag filters, and/or
dust suppression with water along with enclosure of these
operations are adequate to reduce emissions to below 16 kg/hr
(35 Ib/hr). This amounts to about 0.01 grains of particulate
matter per cubic foot (about 23 x 10~6 kg/m3). This is well
below the EPA standard of 0.08 grains/ft3 (183 x 10"6 kg/m3)
for incinerators.
Dust emissions can also occur due to wind, spillage, or
process upsets during conveying operations. Enclosed con-
veyors and dust collectors at transfer points can minimize this
potential problem.
(2) Spent Shale
Disposal methods for spent shale will vary depending on
the type of mining system used.
For a 7. 95 x 106 I/day (50,000 bbl/day) plant, 24.4 to
27.1 x 10^ m tons of shale are required annually. This shale,
having a mined volume of 11. 3 to 12. 7 x 10$ m3, after retorting
has a volume of 17 to 19. 8 x 106 m3. Though compacting can
reduce this volume to 12. 7 to 14. 7 x 10^ m3, the processed
shale still occupies a greater volume than the original mined
shale, so all of it cannot be returned to the mine.
The two most apparent methods for the disposal of the
waste shale are:
Surface disposal of all the spent shale
Return of as much shale as possible to the voids
left by mining with the excess spent shale
sent to surface disposal.
About 10 to 20 percent (by weight) of water is added to the
spent shale to reduce dusting and aid in the consolidation of
disposal piles. Transport to the disposal area may be
accomplished by hooded belt-conveyor or by a water slurry
system.
XII-12
-------
The composition of the mineral content of the spent shale
is shown as Stream _B in Figure XII-1. Some components of
the spent shale are significantly water soluble, indicating the
need to guard against uncontrolled leaching. The natural
hardening reaction, however, aided by moistening and com-
pacting of the spent shale, results in a nearly impervious mass
within a few days. The rapid hardening of the spent shale
should minimize the natural pyrite weathering and resulting
acidic water runnoff.
(3) Wastewater
Wastewater is an inherent by-product of oil shale retorting.
For the size facility analyzed, between 0. 55 x 10^ and 1. 36 x
10^ I/day of retort wastewater can be expected. The upgrading
of the raw shale oil generates about an additional 0. 38 million
liters of wastewater per day. The wastewater, contains a variety
of organic and inorganic components. It is treated with lime,
heated, and contacted with activated carbon and ion exchange
resins. The composition of untreated and treated water is
shown in Table XII-5.
Table XII-5
Untreated and Treated Wastewater
Component
Ammonia
Organic carbon (as complex
mixture of amines, organic
acids, organic bases, and
neutral compounds)
Organic nitrogen
Sodium
Carbonate
Chloride
Nitrate
Sulfate
Grams/Liter
Untreated
Process Water
2.4
2.5
1.0
0.5
20.8
1.8
trace
1.2
Treated
Water
nil
nil
nil
0.06
0.18
0.01
nil
nil
Source: The Shell Oil Company, "The National Energy Position," February 1972, p.22.
XII-13
-------
About 30 x ID**, I/day of wastewater is expected to be
generated from the underground mining operation. Approxi-
mately 50 percent of this water is of poor quality, most of
which is used to wet the spent shale before disposal. The
high quality water (2 to about 7 x 10& I/day) along with 6 x 10^
I/day of poorer quality mine water is disposed of. The dis-
posal of surplus poor quality water requires desalting, evapo-
ration in impermeable ponds, or disposal by subsurface
injection. This is considered part of the mining operation and
does not affect the wastewater treatment requirements dis-
cussed here.
High quality water can be disposed of directly into a river.
In many instances most of the available mine water will be of
poor quality, requiring the high quality water needed for the
plant to be supplied from some other source.
Blowdown from cooling towers and boilers, as well as
sanitary sewage and run-off water, can be treated by a miscel-
laneous treatment section incorporating a sewage treatment
plant. The only waste product discharged from this wastewater
treatment stage will be an inert solid residue.
(4) Sulfur Recovery
Recovery of 64 m tons/day (70 short tons/day) of elemental
sulfur from a stripped H2S stream and from SO2 recovered from
tail gas cleanup is accomplished in a Glaus and Beavon cleanup
unit. It is assumed that 99 percent of the sulfur feed to the
Claus unit can be recovered as elemental sulfur when using this
tail gas cleanup system.
/o\
According to a Bureau of Mines reportv ', H^S is
recovered in the gas treating section by using ammonia-water
as the absorption medium. H^S could have been also recovered
by using a hot carbonate, MEA or DEA process or a solvent-
based system such as an M-Pyrol, Rectisol, or Selexol process.
A Stretford process could be used as an alternative to the
Claus plant.
XII-14
-------
(5) Stack Gas Cleanup
The fuel gas generated by retorting, coking, and hydro-
genating the product fuel, fires the boilers in the steam and
power plant. This gas contains about 1000 ppm of l^S and,
therefore, requires stack gas cleanup following combustion.
For this analysis, stack gas recovery was assumed to be
accomplished by using the Wellman-Lord process. Although
this process has yet to be tested in related applications, it is
expected that this scheme will be demonstrated before the
Gas Combustion Retort process is commercialized.
Several other processes, such as the Bureau of Mines'
Citrate process, limestone injection, lime scrubbing, catalytic
oxidation, and the double alkali process can also be used. A
benefit of using the Wellman-Lord process is that it produces
a concentrated SC"2 stream that may be used to oxidize the
H2S separated from the acid-gas during gas treatment scrubbing.
Alternatively, the SC>2 may be reduced to elemental sulfur by
the Allied process.
A further alternative, resulting in still lower overall
plant emissions, is to use a more effective acid-gas treatment
for the initial gas washing. As was indicated earlier, this
process was demonstrated before the present concern for our
environment. In this analysis (as in the other process synopses)
the process itself was kept intact as designed, and the abatement
processes were simply added. In this case, however, a more
desirable alternative for cleaner emissions is to replace the
ammonia-wash tower system with a more effective acid-gas
treatment. With this substitution, the stack gas cleanup
process would not be required, resulting in overall process
economy.
4. COST OF CONTROL PROCESSES*
The costs of the pollution control and treatment processes de-
scribed in the previous sections of this process synopsis are calcu-
lated here. The processes to extract oil from oil shale are still under
development; thus, the cost figures reported here are based largely
The costs presented here are preliminary estimates; the actual
costs of emission control in this process will not be known until
a full-scale facility has been constructed and operated.
XII-15
-------
on extrapolations from process development unit data and from related
commercial experience. The major assumptions and conventions adopted
are discussed in Chapter III. All by-product flow rate data presented in
Figures XII-1 also apply.
The incremental* capital and operating costs for the control of
potential pollutant and waste streams are presented in Tables XII-6
and XII-7. The incremental investment and operating costs are based
on the control methods selected and described in Sections 2 and 3 of
this chapter. The incremental capital investment of pollution control
facilities using utility financing is $22. 3 million; using the discounted
cash flow (DCF) method, it is $23. 6 million. The annual operating
costs, after taking by-product credits for sulfur at $10/long ton and
ammonia at $25/short ton, is about $1.7 million per year.
The derivation of the formulae for calculating the incremental
cost of shale oil production due to pollution control is presented in
Chapter III. For the DCF method of financial accounting, the required
incremental annual cost of oil, X for the assumed rate of return is:
X = N+ 0.23816 I = 0. 1275 S +0.23077 W
where:
N = incremental net operating cost
I =• incremental plant investment
S - startup costs
W = incremental working capital
For the utility financing cast it is:
X - N + 0. 1198 C + 0.0198
where:
C = required incremental capital investment
N = defined above
W = defined above.
The annual oil production rate is calculated on the basis of a
7. 95 x 106 I/day (50,000 bbl/day) facility. The annual oil production,
G, for this plant is:
* Incremental costs in this report refer to that portion of the total
plant cost that is assignable to pollution control facilities.
XII-16
-------
Table XII-6
Incremental Capital Investment
Item
$ Million
Incremental plant investment
Shale dust emission control
Wastewater treatment
Wellman-Lord SC>2 recovery
Claus sulfur recovery with Beavon tail gas treatment
Subtotal plant investment
Project contingencies0
Incremental plant investment (I)
Startup costs (S)d
Interest during construction6
c
Incremental working capital (W)
Incremental capital investment (C)
5.5
3.2
6.0
1.2
15.9
2.4
18.3
0.6
Utility
3.1
0.3
22.3
18.3
0.6
DCF
4.3
0.4
23.6
Notes:
a.
Incremental plant investment, return on investment during con-
struction, and working capital are treated as capital costs in
year 0 (the year ending with completion of startup operation).
b. Installed costs, including engineering design costs, contractors'
profit, and overhead.
c. Includes costs for unexpected site preparation and hardware
requirements at 15% of plant investment.
d. At 20% of incremental gross operating cost.
e. For the DCF method, computed as the
discount rate x incremental plant invest-
ment for 1.875 years' average construction
period I (l+i)n - 1(1,12)L8T5 = 1.236761 or
0.23676 I additional investment.
For the utility financing method, computed
as the interest rate on debt x incremental
plant investment x 1.875 years.
f. Sum of materials and supplies at 0.9% of
incremental plant investment and net
receivables at 1/24 of annual incremental
revenue.
XII-17
-------
Table XII-7
Incremental Annual Operating Cost
Labor
Dollars
Direct operating labor (3 men/shift x $5/hr
x 8304 shift hrs/man yr)
Maintenance labor (1.5% of I)
Supervisory (15% of direct operating and
maintenance labor)
Incremental labor cost
Administration and general overhead (60% of
incremental labor)
Other direct costs*
Supplies
Operating (30% of direct operating labor)
Maintenance (1.5% of I)
Incremental cost of supplies
Local taxes and insurance (2.7% of I)
Incremental gross operating cost
By-product credits
Sulfur (84.4 short tons/day x 365 days/yr
x 0.9 capacity factor x S10/LT x 2000 LT/
2240 short ton)
Ammonia (138 short tons/day x 365 days/yr
x 0.9 capacity factor x $25/short ton)
Less by-product credits
Incremental net operating cost (N)
124,600
274,500
59,900
37,400
274,500
459,000
275,400
1,500,000
311,900
494,100
247,600
1,133,300
3,040,400
-1,380,400
1,659,500
*This item includes catalysts and chemicals expended and utilities purchased for the pollution
control processes.
XII-18
-------
(50,000 bbl/day) (42 gal/bbl) (6.63 Ib/gal) (19,925 Btu/lb)
(.9 capacity factor) (365 days/yr) = 91.13 x 1Q12 Btu/yr. or
23 x 1012 kcal/yr.
This is regarded as a typical production rate of a commercial
oil shale-to-oilplant.
The incremental cost of oil due to pollution control is
computed as:
Annual Cost of Oil
X
X
Annual Oil Production G
91.13 x 1012 Btu/year
The solution of the above equation for the incremental cost of
oil is shown in Table XII-8. When using the DCF method, the incre-
mental cost of oil, as a result of the investment and operating expenses
for the waste system controls, is 6. 8£/million Btu of oil produced.
Using the utility accounting method, it is 4. 8£ /10° Btu.
Table XII-8
Incremental Cost of Oil Due to
Pollution Control for the Gas
Combustion Retort Process
Accounting
Method
DCF
Utility financing
Incremental Annual Cost of Oil
($/Yr.)
6,186,600
4,336,900
$/106 Btu
0.068
0.048 '
5. REFERENCES
(1) Environmental Statement for the Proposed Prototype
Oil Shale Leasing Program, Volumes I, II, and III,
U.S. Department of Interior, September 1972.
.(2) Development of the Bureau of Mines Gas Combustion
Oil Shale Retorting Process, Bulletin 635, Bureau of
Mines, 1966.
XII-19
-------
XIII. THE LURGI PROCESS (AIRBLOWN GASIFIER)
Although it may be uneconomical to pump a low-Btu gas through
pipelines, this type of gas is a valuable fuel for on-site electric power
generation plants or for industrial use. As one would expect, the
technology to develop a high-Btu or a low-Btu coal gasification
process is similar. The Lurgi gasification process is a commercially
available method to produce a low-Btu fuel gas with a heating value of
1800 to 4000 kcal/m* (200 to 450 Btu/ft3). This gas is suitable for
utility fuel, chemical synthesis or as a base for a substitute natural
gas (SNG) as described in Chapter VI of this report.
The wide range in the heating value of the gas produced depends
on whether air or oxygen is used to support combustion (provide the
heat input, by carbon oxidation in the gasifier). The use of air dilutes
the product gas with nitrogen and reduces its heating value. The basic
Lurgi gasifier is a moving-bed type of reactor which can be considered
to have three temperature zones though which the coal passes as it
falls from the hoppers to the grates at the bottom of the gasifier. The
topmost zone dries and preheats the coal by contact with the hot gases
leaving the reactor. In the middle zone where the coal reaches 600 to
800 degrees C (1100 to 1500 degrees F), devolatilization and gasifica-
tion occur. Steam is used to supply hydrogen for gasification. In the
bottom reaction zone, a combustion zone, the carbon remaining in the
coal char is reacted with air to supply the necessary heat for the endo-
thermic gasification reactions occurring in the middle zone. The
nominal operating pressure in the gasifier is 21 kg/cm2 (300 psia).
The commercial application of this process, as mentioned in
the description of the Lurgi process for manufacture of a high-Btu
gas (Chapter VI) is presently limited to use of noncaking coals and
to a specific range of coal particle size. However, recently com-
pleted development work on a modified Lurgi gasifier indicates that
the Lurgi system might be adapted to accept mildly caking coals as
well as coal fines, although perhaps with a decreased efficiency.
The only airblown Lurgi gasifier installed, built for STEAG at
Lu'nen, West Germany, supplies a low-Btu fuel gas for a combined
cycle power plant. This installation exemplifies a modern commercial
XIII-1
-------
application of this process. Though it has had a protracted startup
(about three years), the STEAG plant is understood to now be opera-
tional. This plant does not, however, contain sulfur removal pro-
cesses.
Due to the general lack of accessible data on the Lurgi process,
the quantitative design information on the pollution control aspects of
the process published by the El Paso Natural Gas Company, as made
public in their FPC filing, has been the primary source of data used
in preparing the process summary.
1. PROCESS DISPLAY
A schematic flow diagram for the production of a low-Btu gas
using the Lurgi process is shown in Figure XIII-1. Stream composi-
tions and flow rates are given in Table XIII-1.
(1) Basis for Analysis
The purpose of this discussion is to evaluate the potential
emissions from a possible low-Btu gas producing process utiliz-
ing a Lurgi gasifier. Emphasis is placed on final disposition of
sulfur and the environmentally satisfactory utilization of process
by-products.
The information contained in this process synopsis was
developed from the open literature. The flow sheet (Fig-
ure XIII-1) is typical of a possible application of the process,
and it indicates the approximate types and quantities of potential
emissions that might be expected from this type of process.
Energy and material balances around the Lurgi gasifier,
discussed in this report, were based upon process design work
completed by the El Paso Natural Gas Company. The analysis
shown is for a Navajo seam coal containing 0.69 percent sulfur
(as received basis). Its composition is shown in Table XIII-2.
The process design with suggested effluent cleanup system
presented in this report is based on a gross production of suffi-
cient low-Btu product gas to manufacture a total of 32.8x 10^ kcal/
day (130 x 10^ Btu/day) of gas. Total production capacity is
XIII-2
-------
FIGURE XIII-1
The Lurgi Process
(Airblown Gasifier)
SULFUR REMOVAL AND RECOVERY
X.
ENERGY BALANCE
CARRIER
COAL
TOTAL ENERGY INPUT
FUEL GAS
NAPHTHA IN FUEL GAS
TAR OIL
PHENOL
AMMONIA
ASH
SULFUP
TAR
OTHER
TOTAL ENERGY OUTPUT
i JO6 kalftw
1800.6
1196.7
74,2
M.Q
18.1
29.7
5.]
133.1
290.0
1600,5
1BOO.S
LEGEND
1 ] SULFUR FLOW RATE (Nmlom/hr
| AIR COOLER
V
®
-------
Table XIII-1
Stream Composition
X
H
Gas Components
(Mole %1
co2
H2S
C2H4
CO
H2
CH4
C^H,-
- 6
°2
N-,
Total
Lb-moles/hr (dry gas)
Gm-molcs/sec
(dry gas)
Flow Rates
lb/hr-(kg/hr)
Water
Coal (MAP)
Ash
Tar
Tar oil
Phenol
Ammonia
Sulfur
Naphtha
Total Ib/hr
(kg/hr)
Stream Numbers
1
133,987
(60,777)
531,910
(241,274)
158,771
(72,019)
824.668
(374,070)
2
20.95
79.05
100.00
36.437.1
4.591
3
513,390
(232.874)
513,390
(232,847)
4
14.83
0.23
0.25
17.45
23.26
5.08
0.38
-
38.52
100.00
73.756.3
9.293
3,322
(1,507)
8,511
(3,861)
11,833
(5,368)
5
14.86
0.01
0.25
17.49
23.31
5.09
0.38
-
38.61
100.00
73.595.0
9.273
3,322
(1,507)
8,511
(3,861)
11,833
(5,368)
6
14.83
0.24
0.26
17.46
23.27
5.07
0.37
-
38.50
100.00
1.071.9
135
7
14.86
0.06
0.26
17.49
23.31
5.07
0.38
-
38.57
100.00
1,070.0
135
8
5,234
(2,374)
5,234
(2,374)
9
8,352
(3,788)
158,771
(72,019)
167,123
(75,807)
10
31,117
(14,115)
31,117
(14,115)
11
17,022
(7,721)
17,022
(7,721)
12
3,949
(1,791)
3,949
(1,741)
13
99.64
0.36
100.00
29.2
4
28,733
(13,033)
7,505
(3,404)
36,238
(16,437)
-------
Table XIII-2
Component Analysis (Dry Coal)
Component
Carbon
Hydrogen
Nitrogen
Sulfur
Oxygen
Trace compounds
Ash
Total
Weight
%
58.73
4.30
1.02
0.82
12.12
0.02
22.99
100.00
co ft Q Q
0.81 x 10 m /hr (28.25 x 10 ft /hr) of 1726 kcal/m (194 Btu/
o
ft0) gas. About 20 Lurgi gasifiers are required for this produc-
tion rate. Approximately 10 percent of the fuel output is re-
quired for necessary steam and electricity generation yielding
a net output of about 30 x 109 kcal/day (117.7 x 1Q9 Btu/day).
This quantity of gas, if burned under boilers, could produce
about 550 MW of electrical power. The net coal-to-gas thermal
efficiency of the plant is determined to be 66.5 percent.
(2) Layout and Symbols
The general direction of process flow shown in Fig-
ure XIII-1 is from the coal supply on the left to the product gas
on the extreme right of the flow design, with residuals and by-
products shown along the bottom. The bold line indicates the
flow of the primary gasification process. The encircled figures
refer to the stream compositions shown in Table XIII-1. The
overall energy balance is shown in Table XIII-3 and in Fig-
ure XIII-1.
XIII-5
-------
Table XIII-3
Energy Balance
Carrier
Input
Coal (824,668 Ib/hr x 8664 Btu/lb)
Output
Fuel gas (27.863 x 106 ft3/hr
x 194Btu/ft3)
Lock gas (.405 x 106 ft3/hr
x 194.2Btu/ft3)
Fuel gas for plant fuel
Net fuel gas
Naphtha in fuel gas (881 1 Ib/hr x 18,400 Btu/lb)
Tar oil (17,022 Ib/hr x 17,300 Btu/lb)
Phenol (3949 Ib/hr x 14,021 Btu/lb)
Ammonia (7505 Ib/hr x 9598 Btu/lb)
Carbon content of ash (8352 Ib/hr x
14,086.8 Btu/lb)
Sulfur (5234 Ib/hr x 3983.4 Btu/lb)
Tar (3 1 , 1 1 7 Ib/hr x 1 6,970 Btu/lb)
Losses*
Energy
x 106 Btu/hr
7144.9
5416.6
78.7
5495.3
- 746.5
4748.8
156.6
294.5
55.4
72.0
117.7
20.8
528.1
1151.0
7144.9
x 106 kcal/hr
1800.5
1365.0
19.8
1384.8
- 188.1
1196.7
39:5
74.2
14.0
18.1
29.7
5.2
133.1
290.0
1800.5
*Includes heat losses through air and water cooling, sensible heat of product, by-products and
waste streams, and other unaccounted losses.
XIII-6
-------
Rhombic-shaped units represent intermediate products
(such as boiler feed water), process needs (such as electric
power and steam), and sources (such as process return water),
for which the distribution is not shown. Nonintegral pollutant
cleanup processes are indicated by sloping rectangles. The
extensive treatment and recovery processes which this symbol
represents are discussed in the process description and pollu-
tion control sections of this process summary. For a thorough
in-depth explanation of each licensed cleanup process mentioned,
the reader is referred to readily available process literature
and the bibliography included in this report. Inverted trapezoids
denote residual storage.
2. PROCESS DESCRIPTION .
The specific steps in the Lurgi process for production of low-
Btu gas are discussed below; the generation and disposition of pollutant
streams are also discussed where appropriate. Thejgross gas produc-
tion rate is about 678 x 1()6 ft3/day of gas containing 194 Btu/ft3.
After consuming about 10 percent of the gross output as plant fuel in
the necessary utilities (e.g., raising of steam), the net gas supply is
about 30 x 109 kcal/day (117 x 109 Btu/day), including the heat content
of naphtha in the gas.
(1) Coal Preparation
Coal feed for the Lurgi gasifier is crushed and dried to a
size range of 0. 3 to 3. 2 cm (1/8 to 1-1/4 inches). Any coal fines
generated during the crushing and drying steps can be washed
and sold.
(2) Gasification
Treated coal is fed to the top of the gasifier by means of
lock hoppers using pressurized product gas for pneumatic trans-
port. The lock hopper gas is recovered (stream 6), desulfurized
and used as plant fuel. A coal lock feeds the coal into the gasi-
fication reactor where it is uniformly spread on top of the moving
bed by a rotating distributor. Normal residence time for the coal
in the gasifier is approximately one hour. A revolving grate at
XIII-7
-------
the base of the gasifier supports the fuel bed, removes the ash
into a lock hopper, and allows introduction of steam (stream 3)
and air (stream 2) to the gasifier. Solids and gases move through
the reactor in a countercurrent fashion. About 20 Lurgi gasifiers
are required for the 32. 8 x 109 kcal/day (130 x 109 Btu/day) pro-
duction.
Char leaving the middle reactor zone is burned with in-
coming air to supply heat required for the endothermic gasifi-
cation reactions in the upper parts of the reactor. Hydrogen
for gasification is supplied by reduction of steam with the carbon
in the coal (2H2O + C -» 2H2 + CO2) in the temperature zone be-
tween 595 and 815 degrees C (1100 and 1500 degrees F). Crude
product gas, after heat exchange with the incoming coal, leaves
the gasifier at 455 degrees C (850 degrees F). In addition to
CC>2> CO, CH4 and H2, the gas leaving the reactor contains
carbonization products such as tar, oil, naphtha, phenols and
ammonia. The gas also contains coal and ash particulates plus
traces of hydrogen sulfide and carbonyl sulfide.
(3) Gas Quench and Cooling
Hot crude gas leaving the gasifier is cooled rapidly by
quenching with a recycle gas liquor spray (an aqueous stream
produced by quenching and cooling hot gases). The gas is further
cooled in a waste heat boiler to produce 2. 1 kg/cm^ (15 psig)
steam, then by air coolers, and finally by heat exchange with
cooling water.
(4) Gas Purification
Removal of the hydrogen sulfide from the Lurgi product
gas will produce a fuel gas containing about 140 ppm sulfur,
primarily in the form of COS. v It is proposed to remove H2S
through use of a pressurized Stretford process operating at
18. 6 kg/cm2 (265 psia). Normal pressure for the Stretford
This COS content is contingent upon use of the low sulfur coal
selected for this analysis.
XIII-8
-------
2
process is 1.8 kg/cm (25 psia) except in some applications in
gas fields. The Stretford absorber solution contains the active
ingredients of Na2CC>3, Na2"VO3, and an organic oxidation
catalyst. The absorption is specific for H2S; however, the
process does not remove the COS or CO2 from the gas. Carbon
dioxide need not be removed from the gas but additional removal
of COS would be desirable so that the combustion products from
burning this gas would contain minimal sulfur emissions. The
concentration of hydrogen sulfide in the treated gas (stream 5)
is expected to be about 10 ppmv, and well below the range where
its contribution of sulfur in the combustion products would be
meaningful.
Regeneration of the Stretford solution with air (2H2S +
O2 ~* 2H2O + S) recovers the sulfur and no tail-gas cleanup sys-
tem is required. Regenerated solution is recycled to the absorber.
(5) Water Supply
Makeup water to the system could be supplied by wells,
surface streams, or reservoirs. Definitive water requirements
could not be determined from the source documents available but
are expected to be about 2000 gpm. Conventional water treating
techniques are used to make high purity boiler feed waters and
other makeup waters for plant needs. Particulate laden water
from the reactor and from other blowdowns is treated to remove
the ash and to recycle the water where possible.
(6) Utility and Power Generation
Air required for the gasifier is compressed in three stages:
to 6.3 kg/cm^ (90 psia) in the first two stages, and 25.3 kg/cm^
(360 psia) in the third stage.
Steam and power required for the production of fuel gas
are generated on site. The specific steam and power require-
ments could not be determined from available sources but are
estimated to be about 0. 68 x 106 kg/hr of steam (1. 5 x 106 Ib/hr)
and 17. 5 MW of electricity for this facility. The steam required
for gasification is identified by stream 3. About 10 percent of
the fuel gas and naphtha produced (stream 5) and the desulfurized
XIII- 9
-------
gas identified as stream 7 (primarily lock hopper gas) is utilized
for the generation of steam and power.
(7) Energy Balance
The energy balance for the process is presented in
Table XIII-3. The overall efficiency, which includes by-product
heat credit (e. g., tar, oil, sulfur, ammonia, etc.), is about
82 percent for the process; the coal-to-fuel gas efficiency is
about 68. 7 percent, including the naphtha, and 66. 5 percent
without it (net). If the tars and oils are hydrotreated to produce
light oils, then the overall efficiency will drop below 82 percent.
(8) Sulfur Balance
The sulfur balance for the Lurgi process flow sheet shown
in Figure XIII-1 is reported in Table XIII-4. The majority of
the sulfur in the coal feed leaves the gasifier as I^S with about
6 percent COS, CS2 and organic sulfur such as thiophene. Most
of that sulfur is recovered as elemental sulfur in the Stretford
unit where sulfur recovery is 93. 8 percent of the sulfur fed to
the units (92 percent of the sulfur in the sized coal feed). Ap-
proximately two percent of the sulfur fed with the coal remains
with the tar and tar oil products. Less than 0. 8 percent of the
total sulfur feed is discharged to the atmosphere in the fuel sup-
plied to the plant utilities; another 0. 2 percent of the feed sulfur
is discharged with a CO2~rich gas (low pressure Stretford II
absorber off-gas containing 0. 12 m ton/day sulfur).
3. DISCUSSION OF POLLUTION CONTROL PROCESSES
There are two approaches to define the pollution control pro-
cesses in low-Btu systems such as the one described in this chapter.
One approach considers the entire gasification process — from coal
handling through gas purification—as pollution control. The rationale
used is that the primary reason for installing the process is to con-
sume coal in a more environmentally satisfactory manner. This ap-
proach was used in the process synopses of this report where direct
combustion or indirect combustion (low-Btu gasification and desul-
furization) of coal or char were used.
XIII-10
-------
Table XIII-4
Sulfur Balance
(29. 7 x 109 kcal/day Plant - Net)
(117.7 x 109 Btu/day)
Carrier
Input
Coal
Output
Fuel gas
Tar
Tar oil
20% Ammonia solution
Elemental sulfur (recovered
from Stretford I sulfur plant)
Sulfur discharged to atmosphere
from required utilities
Sulfur compounds discharged to atmosphere
from Stretford II plant
Sulfur
(Lb/Hr)
5691.5
288.4
84.0
25.5
3.8
5233.3
45.3
11.2
5691.5
(Kg/Hr)
2581.7
130.8
38.1
11.6
1.7
2373.8
20.6
5.1
2581.7
Sulfur
Distri-
bution
%
100.00
5.07
1.47
0.45
0.07
91.95
0.79
0.20
100.00
Note:
Sulfur dioxide emitted from combustion of product fuel gas:
-,oo^,u/u 64.066 lbSO2
288.4 Ib/hr x *•
32.066 Ib S ^ 0.1171bSO2/106Btuor
4905.4 million Btu/hr of ~ (0.211 kg SO2/106 kcal)
fuel gas and naphtha
The alternative approach to define which are pollution control
portions of low-Btu gas manufacturing schemes is based on the
assumption that the product gas is a superior fuel to the original coal
feed and would be manufactured for that reason. As high-temperature
gas turbines are developed for use in combined cycle power genera-
tion, this argument becomes more valid. In this case, gasification
permits utilization of coal in more efficient power plants. Using this
definition, the pollution control processes are considered to include
XIII-11
-------
only those unit processes performing sulfur removal and recovery
functions, wastewater treatment, ash disposition, and recovery of
fines generated in the coal crushing circuit. This approach was used
in this synopsis.
The major waste streams generated in the low-Btu Lurgi gasifi-
cation scheme of Figure XIII-1 are presented in Table XIII-5. The
sources of the potential pollutants and their treatment to minimize
emissions from the process are also presented in that table and are
further discussed here.
Foul water streams known as gas liquors are separated from
the primary gas process stream by condensation during cooling. The
primary difference between the tarry gas liquor and the oily gas liquor
is the temperature at which they are condensed. The higher tempera-
ture condensate (the tarry gas liquor) is condensed with the high-
molecular-weight tars from the quench tank and waste heat boiler
downstream of the gasifier. The tars are decanted, recovered and
sold for their by-product value. The tarry gas liquor is recycled,
after cooling, as scrub liquor for the quench tank, and a portion is
withdrawn to the tar-oil separator where it blends with the oily gas
liquor. The oily gas liquor condenses out of the gas stream with the
lower-molecular-weight tar oils during air and water cooling. While
the tar oils are recovered for their by-product value (stream 11), the
oily gas liquor proceeds to wastewater treatment. The systems to
this point are not considered "emission control" because cooling is
required in the basic process; the tars and oils would normally be
recovered for their by-product value regardless of environmental
considerations. However, the gas liquors from this operation must
be further treated because of the phenol, sulfur, ammonia, cyanide,
and other contaminants present.
The flash gases from the gas liquor expansion vessels will con-
tain hydrogen sulfide, carbon dioxide, as well as some fuel gases that
were dissolved in the liquor. These fuel gases are sent to sulfur re-
covery (discussed later).
The gas liquors, after expansion for flash gas recovery and de-
cantation for tar and tar oil recovery, are transferred to the phenol
recovery section. Here, the gas liquor is sol vent-extracted in a
Lurgi Phenosolvan process to recover the phenols for their by-
product value (stream 12).
The dephenolated gas liquor, after solvent stripping, passes
through a deacidifier where acid gases such as ^S and HCN are re-
moved. The deacidified acid gases are sent to the low pressure
XIII-12
-------
Table XIII-5
Wastes, Their Amounts, Sources, and Treatments for a Coal to Low-Btu Fuel Gas Plant
Final Wastes
Coal fines
Ash
Gas liquor
Hydrogen sulfide,
carbonyl sulfide
and carbon
disulfide
Sulfur dioxide
Amounts
75, 807 kg/hr
( 167.T23 Ib/hr)
31,942kg/hr
( 59, 593 Ib/hr)
tar, tar oil,
phenols, ammonia,
H2S, etc.
2462.4 kg/hr*
(5521.71b/hr)*
About 20.9 kg/hr
( 45 Ib/hr) of
sulfur
Sources
Coal feed, crushing
handling, etc.
From gasifiers
Quenching and gas
cooling system
With the fuel gas
Gas-fired turbines,
boilers, heaters,
incinerators
Treatment
Cyclone separators, bag filters,
scrubbing or washing, etc.
Sluice way and return to mine
for land fill.
Using gas liquor separation sys-
tem, phenol extraction system
and gas stripping system.
Sulfur is recovered using the
pressure Stretford process.
COS and CS? are not recovered;
5.3 kg/hr ( 1 f.7 Ib/hr) of sulfur is
emitted to the atmosphere.
None
*Intermediate waste stream
-------
Stretford (II) unit for sulfur recovery, cyanide conversion and are
then vented (off-gas) to the atmosphere. The deacidified gas liquor
is next steam-stripped to recover a by-product ammonia solution for
sale (stream 13). The treated gas liquor should now contain only small
ppm quantities of phenols, cyanides, and sulfur and should be a satis-
factory input to the plant cooling water circuit where the remaining
impurities would be destroyed by natural biological oxidation.
The ash collected from the gasifiers should be low in carbon
content and contain almost none of the original sulfur from the feed-
stock. That sulfur which is still present is firmly bound into the
organic lattice of the remaining carbon. The ash is quenched with
recycled water from the waste treatment area and slurried through a
system of solids dewatering equipment. The water is recycled to
wastewater treatment and the ash is returned to the mine for disposal
(stream 9).
The Lurgi gasifier will not accept coal fines less than 2 mm in
size. In other process synopses in this report, fines were generally
minus 200 mesh and became airborne during coal grinding. In this
case, however, the majority of the fines are much larger (and could,
in fact, be used as a feed in most other process gasifiers). As much
as 25 to 40 percent of the initial, "as mined" coal will fall into this
size range and must be disposed of (e. g., sold) or briquetted with
coal tar and then gasified.
In the process described here, sulfur removal from the main
process gas stream arid elemental sulfur recovery are achieved in
the same unit process - a high-pressure Stretford system. In most
of the other processes described in this report, the sulfur (and per-
haps the carbon dioxide) were removed from the process gas stream
by an acid-gas treatment system and then the sulfur was separated in
a recovery system. In this Lurgi low-Btu gas process, however,
complete removal of the sulfur is not required and removal of carbon
dioxide was not desired (some of this process gas will be used in gas
turbines to raise power and the carbon dioxide represents mass at
pressure that can be used to drive turbines and generate power).
The Stretford system removes hydrogen sulfide, hydrogen
cyanide (if hydrogen sulfide is present), but not carbonyl sulfide,
carbon disulfide, or other organic sulfur compounds. Similarly, car-
bon dioxide is not attacked as in many other sulfur removal processes.
XIII-14
-------
Of the total sulfur fed to the facility in the sized coal (2581. 7 kg/
hr), 51. 4 kg/hr appears in the liquid by-product streams and 2530. 3
kg/hr or 98 percent appears in the gases. Of this quantity, 2373. 8 kg/
hr is recovered as elemental sulfur and 156. 5 kg or 6.2 percent of the
gaseous sulfur is not recovered by the Stretford system and remains
in the gas as COS, C^S, etc. The concentration of this unrecovered
sulfur in the fuel gas is 139 ppmv.
Of the total quantity of fuel gas manufactured, a significant
quantity exists at low pressure from lock gases (stream 6). The lock
gas has a heating value similar to the primary gas stream and it should
be desulfurized for use. Rather than recompress this gas back to the
gasifier system pressure, one of the absorbers in the Stretford sys-
tem was operated at low pressure (low-pressure Stratford I absorber).
Similarly, the deacidification off-gas and the expansion gases from
the gas liquors contain sulfur that should be removed before these
gases are vented. Another low-pressure Stretford absorber is used
to treat these gases (low-pressure Stretford II absorber). The com-
bined rich liquor from all three Stretford absorbers in this process
are oxidized for sulfur recovery in a single unit (stream 8).
Most of the effluent from the high-pressure Stretford unit makes
up the low-Btu product gas. However, a small side stream is routed
to the low-pressure fuel product (stream 7) to supply the energy re-
quirements for the utilities of the facility. Approximately 10 percent
of the total energy output is used for this purpose. The total sulfur
emitted to the atmosphere during on-site combustion of this captive
utility fuel gas is 90. 6 Ibs SO2 per hour (41. 1 kg/hr) or 0. 117 Ibs SO2
per million Btu consumed within the process (0. 211 kg/106 kcal).
The net plant gas output, when consumed, will emit 576. 2 Ibs SO2
per hour to the atmosphere (261.4 kg/hr) or about 0.211 kg SO2/million
kcal (0. 117 Ibs SO2 per million Btu). This level of sulfur emission, of
course, is related to the low sulfur content of the coal feeding this fa-
cility. If a higher sulfur feed (e. g., Eastern coal) were used, the low-
Btu product gas may contain up to 900 ppm sulfur (0. 75 Ibs SO2/10^
Btu or 1. 35 kg/10^ kcal). These emissions, even with high sulfur coal
feedstock, compare favorably to the present Federal standards for
direct combustion of coal of 2. 16 kg SO2/106 kcal of fuel burned
(1.2 Ibs SO2/1Q6 Btu).
XIII-15
-------
4. COSTS OF POLLUTION CONTROL*
The costs of pollution control are dependent on which of the two
definitions are used for pollution control facilities. Under the first
assumption discussed in the section above, the entire process is con-
sidered as pollution control since it permits the combustion of coal.
Under the second assumption discussed above, only specific op-
erations within the facility are considered as pollution control. The
capital investment of these pollution control and sulfur treatment
processes was based on an economic evaluation made by Stearns -
Rogers, Inc. , for the El Paso Natural Gas Company's SNG Lurgi
plant (reference 1). The investment and operating costs for the gas
liquor separation, extraction and stripping sections, as well as the
sulfur recovery section, are based on the amount of coal going to the
airblown gasifier in the El Paso plant. For a Lurgi plant producing
it it
130 billion Btu/day (gross) of utility gas, the incremental invest-
ment costs for the pollution control and treatment process are given
in Table XIII-6 and the detailed operating costs are given in Table XIII-7.
For the discounted cash flow method (DCF), the required incre-
mental annual cost of gas, X, for the assumed rate of return is:
X = N + 0.238161 + 0. 1275S + 0. 23077W
where
N - Incremental net operating cost
I - Incremental plant investment
S = Startup cost
W = Incremental working capital.
* The costs presented here are preliminary estimates; the actual
costs of emission control in this process will not be known until
a full-scale facility has been constructed and operated.
** Incremental costs in this report refer to that portion of the total
plant cost that is assignable to pollution control facilities.
XIII-16
-------
Table XIII-6
Incremental Capital Investment ($1000)
(117. 7 Billion Btu/Day Net Plant)
Item
Fuel gas treating
Gas liquor separation
Phenol extraction
Gas liquor stripping
Sulfur recovery
Subtotal plant investment
Project contingencies
Incremental plant investment (I)
Startup costs (S)
Interest during construction
Incremental working capital (W)
Incremental capital investment
$ Thousand
6,188
2,682
3,349
2,918
3,916
19,053
2,858
21,911
533
Utility
3,697
350
26,491
21,911
533
DCF
5,188
439
28,071
For the utility financing case, it is
X = N + 0. 1198C + 0.0198W
where
C = Required incremental capital investment
N = Incremental net operating cost
W = Incremental working capital
XIII-17
-------
Table XIII-7
Incremental Annual Operating Cost ($)
(117. 7 Billion Btu/day (Net), 90% Stream Factor)
Item
Cost ($)
Direct operating labor (6 men/shift x
$5/hr x 8,304 hr/year
Maintenance labor (1 .5% of I)
Supervision ( 1 5% of direct operating
and maintenance labor)
Incremental labor cost
Administration and general overhead
(60% of incremental labor)
Other direct costs
Supplies
Operating (30% of direct operating labor)
Maintenance ( 1 .5% of I)
Incremental cost of supplies
Local taxes and insurance (2.7% of I)
Incremental gross operating cost
By-product credits
Crude phenol (3949 Ib/hr x 24 x 328.5
x $ .04/lb)
Sulfur (1234 L1yhr x 24 x 328-5 x $io/LT)
Ammonia
ton/hr x 24 x 323.5 x
$25/ton)
Operating Cost
249,000
329,000
87,000
665,000
399,000
1,233,000
75,000
329,000
404,000
592,000
1,245,000
184,000
740,000
2,169,000
665,000
399,000
603,000
404,000
592,000
2,663,000
(2,169,000)*
494,000
* By-product credits are not allowed for tars and tar oils since these are recovered
during gas cooling (defined as a requirement of the basic process scheme - not a
pollution control). Simiarly coal fines are screened out as a process requirement
and therefore by-product credits obtained from their sale should not be used to
effect the cost of pollution control.
XIII-18
-------
Annual gas production
G = (474.8 x 156. 6) x 106 Btu/hrx 24 x 328.5
= 38,674,174 x 106 Btu/year
Therefore, the incremental cost of gas due to pollution control is
X _ Annual incremental cost of gas
G Annual gas production
The incremental cost of gas due to pollution control is shown in
Table XIII-8. These costs do not include by-product credit for tar
and tar oil. However, if these by-product credits are taken, then the
costs given will be reduced from 15.2£/106 Btu to 6. 7
-------
XIV. THE U-GAS PROCESS
The U-Gas process is a system being developed by the Institute
of Gas Technology (IGT) for the fluid-bed gasification of coal with
steam and air (or oxygen), to produce a low-Btu utility fuel gas. Ad-
vanced R&D and pilot plant scale studies for development of the U-
Gas concept are the subject of proposals now pending before various
agencies.
The U-Gas process will produce a gas with heating value rang-
ing from about 150 to 300 Btu/ft3* (1335 to 2670 kcal/m"), depending
on whether air or oxygen is used as the gasifying reactant. »
Preliminary designs have been developed for a plant which will
produce a low-Btu fuel gas to serve an electrical generating plant for
use in two possible applications:
Production of low-Btu gas to use as feed in a combined
cycle power station, anticipating that high-temperature
gas turbines for use in a combined cycle operation will be
commercially available in the next ten years. (The flow
sheet for this process is given in Figure XIV-1.) For
this application an airblown gasifier may be employed be-
cause nitrogen dilution is not detrimental. However, the
gas must be delivered at elevated pressure, requiring
high pressure desulfurization. Ideally, the water vapor
and carbon dioxide in the raw gas should not be removed
because they represent mass, at elevated temperature
and pressure, that may be utilized as motive power to
help drive the turbines.
Production of low-Btu gas to be used as fuel for large in-
dustrial users or conventional electrical generating facil-
ities. (The flow diagram for this application is given in
* Throughout this synopsis, ft3 refers to one cubic foot when
measured at standard conditions of 60 degrees F, 30" Hg and
saturated with water.
XIV-1
-------
FIGURE XIV-1
The U-Gas Process—Combined Cycle Case
X
<
to
HEAT TO STEAM CYCLE
157.8 x 106 kcal/hr
o
PRETREATEH GAS
EXTERNAL
CYCLONE
SEPARATORS! 20-7 k9'
CYCLONE
SEPARATO
PRETREATER
21.1 kg/cm2
425°C
PITTSBURGH SEAM \rotor
19.2 x 106 kcal/hr
SATURATED STEAM
FROM PRETREATER
100X 106
kcal/hr
ST6AM
TO BFW
SUPEP
HEATER / WATER-FILLED
HOPPER
WHB- Waste Heal Boile
BFW . Boiler Feed Wate
A
FRJOM
COMPRESSOR
STEAM TO SELEXOL
HEAT FROM WHB-1
157.8x 106kcal/h
SELEXOL
ACID-GAS
REMOVAL
CHEVRON
WASTEWATER
TREATMENT
CLAUS
SULFUR RECOVERY
RECYCLE TO
COOLING TOWERS
(CONTAINS 0.01 8">-mol/sec
OF SULFUR
AIR COMPRESSOR -—
132.000 Hp INTERSTAGE
0
ENERGY BALANCE
CARRIER
COAL TO PROCESS
TOTAL INPUT
PRODUCT GAS
GAS EXPANSION ENERGY
BY-PRODUCT STEAM
SULFUR
AMMONIA
ASH
LOSSES
PROCESS REOMTS
TOTAL OUTPUT
x,06
1911.0
1398.9
6.0
227.8
24.3
0.4
84.5
276.5
-107.4
kcal/hr
1911.0
1911.0
•ENERGY SUPPLIED FROM COMBINEDCYCLE. OTHER
PLANT POWER REQUIREMENTS INCLUDING PUMPS,
TURBINES. FANS. ETC., UTILIZING ENERGY FROM
COMBINED CYCLE AMOUNT TO 25,593 kWh/hr.
BY-PRODUCT
STEAM
227.8 x 106 kcal/hr
CONTAINS 2 gnvmol/sec
OF SULFUR
-------
Figure XIV-2.) An oxygen-blown gasifier is employed
in this case to provide a richer gas. High pressure is
not required, so the gaseous product may be desulfurized
at near-ambient pressure conditions.
1. PROCESS DISPLAY
The schematic flow diagrams and energy balances for the two
applications of the U-Gas process described in this chapter are shown
in Figures XIV-1 and XIV-2. The stream compositions and flow rates
are presented in Tables XIV-1 and XIV-2.
(1) Basis for Analysis
The process flow sheet data, stream compositions and flow
rate data presented in this synopsis were modified from input
developed by the process developer, IGT.
The two applications of the U-Gas process described here
include the production of a low-Btu gas in an oxygen-blown gasi-
fier to be used as fuel by industrial users or to raise steam in
a conventional steam power plant (the fuel gas case) and the
production of a low-Btu gas in an airblown gasifier for use in a
combined cycle power station (the combined cycle case). In
both cases, the feed for the process is 7346 tons/day (6664 m tons/
day) of coal (as received at 6 percent moisture by weight). This
amount is sufficient to generate approximately 1000 MW (gross*)
of electrical power in the combined cycle scheme. Data are
based on the use of Pittsburgh Seam coal, which contains 4. 4 per-
cent sulfur (dry basis) and has a heating value of 12,388 Btu/lb
or 6882 kcal/kg. The composition of this coal is given in
Table XIV-3.
* The net power output of this facility would be about 5 percent
lower due to energy requirements fed back to the combined
cycle system.
XIV-3
-------
FIGURE XIV-2
The U-Gas Process—Fuel Gas Case
RAW
X
t—(
t
7179 gm mof/sec
AIR | COMPRESSED AIR,
TO PRETREATER
POWER RECOVERY SUMMARY
POWER FROM GENERATOR 1
POWER FROM GENERATOR II
TOTAL GENERATED
LESS PLANT MOTORS
NET BY-PRODUCT POWER
fcWh
73.900
25.340
99.240
-26,670
72,570
ATMOSPERE
CONTAINS 2 gm-mol/sec
OF SULFUR
I
ENERGY BALANCE
CARRIER
COAL TO PROCESS
TOTAL INPUT
PRODUCT FUEL GAS
BY-PRODUCT POWER
SULFUR
ASH
LOSSES
TOTAL OUTPUT
x 106 kcal/hr
1911.0
1515.9
62.4
24.1
78.5
230.1
1911.0 .
1911.0
-------
Table XIV-1
Stream Composition—Combined Cycle Case
X
t—H
Stream Description
Stream Number
kg/I"
Component
C
H
O
N
S
Ash
Coal Feed
A
277,667
wt 7, (dry)
71.5
5.00
6.50
1.20
4.40
11.4
100.00
Pretreated Char
B
238,296
wl Si (dry)
71.25
4.02
7.50
1.00
3.74
12.49
100.00
Ash
C
39,212
wt 5 (dry)
20.33
1.43
-
1.78
0.58
75.88
100.00
Strain Description
Stream Number
Temperature, °C
Pressure, kg/cm
Component
CO
CO2
H2
H2O
CH4
Ni
NH3
HCN
H2S
COS
S02
C2«6
Tar
Total, ItMnol/hr
(gm-mol/sec)
Pretieater
Gas
1
427
21.1
mot -7,
2.91
7.97
23.02
0.46
64.66
0.1
a 25
0.03
loaoo
Ib/mol/hr
735
2,011
5,806
115
16,311
176
63
8
25,225
(3,178)
Raw Gas
2
843
21.1
molS
16.97
8.77
11.58
12.00
4.12
45.85
0.01
0.68
0.02
100.00
Ib/mol/hr
18,595
9,609
12,686
13,148
4,516
50,246
9
0.3
750
24
109,583.3
(13,807)
Water to Waste
Water Treatment
3
mol%
0.02
0.06
0.04
99.77
0.02
0
0.07
0.02
100.00
Ib/mol/hr
2
8
5
12,820
2 '
9
0.3
2
12,848.3
(1,619)
Water Effluent
4
38
moll
0.05
99.95
100.00
Ib/mol/hr
6
12,820
Sppm
12 ppm
10 ppm
12,826
(1,619)
To
Sulfur Recovery
5
38
1.4
mol%
15.04
15.04
37.58
15.04
2.26
15.04
100.00
Ib/mol/hr
-
2
5
...
2
...
0.3
2
13.3
(1.7)
Gas to Setexol
6
38
20.0
mol%
19.22
9.93
13.11
0.34
4.67
51.94
0.77
0.02
100.00
Ib/mol/hr
18,593
9,601
12,681
32S
4,514
50,246
748
24
96,735
(12,189)
Selexol Effluent
7
38
19.7
mol%
20.16
6.72
13.75
4.89
54.47
0.005
0.01
100.00
Ib/mol/hr
18,593
6,198
12,681
...
4,513
50,246
5
12
92,248
(11,623)
H2S
OfFGas
8
93
1.4
mol%
75.84
7.31
0.02
16.56
0.27
100.00
Ib/mol/hr
...
3,403
...
328
1
...
743
12
4,487
(565)
-------
Table XIV-2
Stream. Compositions—Fuel Gas Case
X
OS
Stream
Stream Name
Flow Rate (kg/hr)
Component
C
H
O
N
S
Ash
Total
A
Coal Feed
277,667
wt % (Dry)
71.50
5.00
6.50
1.20
4.40
11.40
100.00
B
Pretreated Char
238,296
wt % (Dry)
71.25
4.02
7.50
1.00
3.74
12.49
100.00
C
Ash
38,924
wt % (Dry)
20.06
1.09
-
1.80
0.59
76.46
100.00
Stream
Stream Name
Pressure (kg/cm^)
Temperature (°C)
Component
CO
COi
H2
H2O
CH4
N2
H2S
SO2
C2H6
Tar
COS
HCN
NH3
Total
1
Pretreater Gas
22
425
mol%
2.91
7.97
-
23.02
0.46
64.66
-
0.70
0.25
0.03
-
-
-
100.00
gm-mol/sec
93
253
-
732
15
2055
-
22
8
1
-
-
-
3178
2
Raw Gas
22
770
mol%
25.59
11.08
17.66
15.43
7.21
21.99
1.00
-
-
-
0.03
-
0.01
100.00
gm-mol/sec
2410
1043
1663
1452
679
2071
95
-
-
-
3
0.04
1
9417
3
Fuel Gas
2.4
65
mol%
25.85
11.19
17.84
15.58
7.28
22.22
10 ppmv
-
-
-
0.03
-
0.01
100.00
gm-mol/sec
2410
1043
1663
1452
679
2071
< 1
-
-
-
3
-
1
9322
-------
Table XIV-3
Component Analysis (Dry Coal)*
Component
Carbon
Hydrogen
Oxygen
Nitrogen
Sulfur
Ash
Total
Weight %
71.50
5.00
6.50
1.20
4.40
11.40
100.00
This coal has caking characteris-
tics and requires pretreatment.
(2) Layout and Symbols
The general direction of process flow is from the coal sup-
ply on the left to the product gas on the extreme right of the flow
diagram, with residuals and by-products shown along the bottom.
The bold line indicates the flow of the primary gasification pro-
cess. The encircled figures and letters refer to the stream com-
positions shown in Tables XIV-1 and XIV-2. The overall energy
balances are discussed later in this chapter.
Rhombic-shaped units represent intermediate sources
(such as process waste heat generation) for which the distribu-
tion is not shown.Nonintegral pollutant cleanup processes are
indicated by sloping rectangles. The extensive treatment and
recovery processes which this symbol represents are discussed
in the process description and pollution control sections of this
process summary. Inverted trapezoids denote, residual storage.
XIV-7
-------
2. PROCESS DESCRIPTION
The unit processes which comprise the gasification phase of the
U-Gas system are described in this section. Further discussion of the
gas purification and pollution control phases of the process is discussed
in Section 3 of this synopsis.
(1) Coal Preparation and Pretreatment
In the U-Gas process, coal of any rank—from bituminous
to lignite — is first sized and pressurized in a lock hopper. If
the coal has caking tendencies, it is transferred to a pretreater
where it is'reacted with air at 700 to 800 degrees F (370 to 425
degrees C) and 300 psia (21 kg/cm^) to oxidize a small portion
of the coal and render it noncaking. The coal residence time in
this pretreatment reactor is about 30 minutes. The air require-
ment is approximately one pound of air per pound of coal. In the
design used for this analysis, the heat of reaction, 520 million
Btu/hr (131 x 1Q6 kcal/hr), is recovered by coils immersed in
the fluidized bed. All the steam required for subsequent gasifi-
cation is generated with this heat. Some excess thermal energy
is available to raise steam for generating by-product power
[discussed later in Sections (5) and (8)] .
(2) Gasification
The U-Gas process utilizes fluidized bed gasification. The
advantages of this type of operation include the following:
High reaction rates can be attained
Coal fines from mining and crushing can be used in
the feed
The mass of carbon in the fluid bed ensures reduc-
ing conditions at all times
The bed temperature can be readily controlled.
Treated coal overflows the pretreater into the fluidized
bed gasification reactor, which also operates at 300 psia
XIV-8
-------
(21 kg/cm.2). The effluent gas from the pretreater stream 1
discharges into and mixes with the gasifier overhead stream.
1. Combined Cycle Case
In the combined cycle case, the pretreated coal is
reacted with a mixture of air and steam, using approxi-
mately two pounds (0. 9 kg) of air* and 0. 5 pounds (0.2 kg)
of steam for each pound (0. 5 kg) of initial coal fed to the
system. Gasification takes place under nonslagging con-
ditions at about 1900 degrees F (1040 degrees C) in the
fluidized bed. The gasifier operates with an average gas
velocity of 1 ft/sec (0.3 m/sec) and a char residence time
of about 50 minutes. When using airblown gasifiers, five
reactors (plus one spare), each with an inner diameter of
22. 5 feet (6.9 m) and a shell length of 30 feet (9. 1 m), are
required for the 7346 tons/day (6664 m tons/day) of coal
feed.
To protect the discharge lock hoppers from the hot
ash agglomerates, they are filled with water to absorb the
heat contained in the ash. When filled, the hopper contents
are flushed into separators to recover ash for final dis-
posal.
The gasifier design includes a free space above the
fluidized bed, giving a gas residence time of 10 to 15 sec-
onds; the gas here reaches a temperature of between
. 1500 degrees - 1900 degrees F (815 degrees - 1040 de-
grees C). With this design, the liquid hydrocarbons that
may evolve are thermally hydrocracked to gas: no tars,
oils, or phenols are expected in the raw gasifier product.
Elimination of tars and oils from the raw gas (stream 2)
reduces heat-exchanger fouling and simplifies waste
stream treatment.
The power to drive the air compressor (131, 900 hp) would be
taken from the shaft of the gas turbine in the power generation
section of the facility.
XIV-9
-------
Most of the dust contained in the effluent gas is re-
moved by internal cyclone separators and returned directly
to the fluidized bed. Fine dust is separated in an external
cyclone, a second stage of dust removal, and is injected
back to the gasifier where carbon contained in the fine
dust is gasified. The fine ash sticks to the heavy agglom-
erates and is removed from the system. After fine dust
recovery, the combined pretreater and gasifier off-gases,
with a heating value of 152 Btu/ft3 or 1353 kcal/m3(dry
basis), leave this section at 1550 degrees F (843 degrees C).
They are then cooled to 66 degrees C by two waste heat
boilers and air coolers.
2. Fuel Gas Case
For the fuel gas production case, the use of oxygen
instead of air as the coal-gasifying reactant raises the
heating value of the raw gas to 255 Btu/ft3 (2269 kcal/m3),
or on a wet basis, 214. 8 Btu/ft3 (1911 kcal/m3). In this
case, the pretreater is operated with air, and 99. 5 per-
cent purity oxygen is used in the gasifier. The same se-
quence of operations is used as for the combined cycle case,
except that an oxygen facility with a capacity of 2650 tons/
day (2400 m tons/day) has been added to the plant. This
facility is powered by shaft energy from the expansion tur-
bine located downstream in the process. First, air is
compressed to 100 psia (7 kg/cm2) and cooled; then 37 per-
cent is further compressed to 350 psia (24.6 kg/cm2) with
expander shaft power and delivered to the pretreater. The
remaining 63 percent goes to oxygen separation. The re-
sulting high purity oxygen is compressed to 350 psia
(24. 6 kg/cm2), also with expansion turbine shaft power.
The total power required for this series of compressions
is 90, 600 hp, compared with 131, 800 hp for airblown gasi-
fication. Of course, air compression for the pretreatment
and the oxygen plant could have been supplied from external
sources. In this case the power output for the expansion
turbine would have been increased. This alternative was
not considered here because of the ready availability of
shaft energy and the less expensive option as presented.
XIV-10
-------
With oxygen-blown gasification, the total raw gas
(gasifier plus pretreater effluent) is 68 percent of that with
airblown gasification, so the number of gasifiers can be
reduced from six to four (one spare in each case).
(3) General Comments
The oxygen-blown gasifier generates product fuel gas at a
rate of 144 billion Btu/day (36.3 x 109 kcal/day) from the specified
coal feed, compared with 129 billion Btu/day (32. 5 x 109 kcal/day)
when using an airblown gasifier system. In both cases, a sub-
stantial amount of additional heat can be recovered and used to
generate power in a steam cycle. For the airblown gasifier,
21. 7 billion Btu/day (5.5 x 10" kcal/day) is recovered in this
manner; about half this amount is recovered from the oxygen-
blown system. Therefore, the total energy recovered is nearly
the same in both cases, about 150 billion Btu per day (37. 8 x 109
kcal/day).
At this point in the process, the flow schemes differ de-
pending on the final use of the gas and the sulfur recovery system
used.
(4) Gas Purification, Combined Cycle Case
This section will include a discussion of the waste heat
recovery, sulfur removal and wastewater treatment portions of
the U-Gas process.
In the combined cycle case, the product gas is required at
pressure to provide turbine motive power. The gasifier is air-
blown since nitrogen dilution is acceptable. Therefore, a sulfur
removal process should be selected that will operate at the gasi-
fier pressure. The Selexol process was chosen as the acid-gas
treatment for this analysis. This acid-gas removal process
requires a low-temperature feed (relative to gasifier raw gas
temperature). Equipment must be provided to cool the gas and
to process the condensed water.
The 800 degree F (425 degree C) gas is further cooled in
a second waste heat boiler to 275 degrees F (135 degrees C).
XIV-11
-------
This recovered heat is used to generate steam required in the
Selexol unit. The gas is then cooled to 150 degrees F (66 de-
grees C) in an air cooler, and final cooling is accomplished in
a scrubber—a direct contact device that scrubs ammonia
(stream 3) from the gas stream. The cool, scrubbed gas
(stream 6) is then passed through the Selexol process (discussed
later) for removal of most of the sulfur. Some carbon dioxide is
also removed in this process, increasing its heating value to
159 Btu/ft3 (1415 kcal/m3). The gas can now be used as a clean
fuel in a combined cycle power facility.
(5) Energy Balance, Combined Cycle Case
The overall energy balance for the process is presented
in Table XIV-4.
The gasifier-power system should be best analyzed as a
unit, because of the energy transfer between them. Hence, the
energy balance for the U-Gas gasification system designed for
use in combined cycle operation is difficult to specify unambigu-
ously. For example, the air compressor for the gasifier would
be driven by shaft horsepower from the gas turbine of the com-
bined cycle system. Similarly there is compression energy in
the process gas. Therefore, the energy balance of Table XIV-4
is based on the total net energy transferred with an overall coal-
to-energy efficiency of 79. 8 percent.
At this stage in the development of the process, the over^
all steam, water and electrical requirements of the gasifier
portion of the system (as described in this analysis, with low
temperature desulfurization) have not been estimated.
(6) Sulfur Balance, Combined Cycle Case
The sulfur balance for this process, which uses 4. 4 per-
cent sulfur (dry basis) Pittsburgh Seam Coal, is shown in
Table XIV-5.
This process recovers about 96 percent of the sulfur in
the coal as a liquid by-product by sending the Selexol off-gas
to a Claus plant with a Wellman-Lord tail-gas scrubber and
XIV-12
-------
Table XIV-4
Energy Balance Calculations, Combined Cycle Case
Carrier
Coal
Total energy input
Product gas
Net expansion energy in product gas*
By-product steam
Less: Air compression requirements
Electric power requirements
Net energy outputt
Sulfur
Ammonia
Ash
Losses
Total energy output
Calculations
612,140 Ib/hrx 12,388 Btu/lb
92,248 mol/hrx 378.5 ft3/mol x 159 Btu/ft3
645,700 Ib/hr x 1400 Btu/lb @ 2400 psig, 1000° F
1 3 1 ,800 hp (heat equivalent)
26,500 kWh@ 341 3 Btu/kWh
7557 mol/hr x 32.06 Ib/mol x 3983 Btu/lb
9 mol/hr x 1 7.03 Ib/mol x 9598 Btu/lb
86,446 Ib/hrx 3878 Btu/lb
(By difference)
Energy
(x!06Btu/hr)
7583.2
7583.2
5551.2
23.8
904
-335.6
-90.4
6053
96.5
1.5
335.2
1097
7583.2
(x!06kcal/hr)
1911.0
1911.0
1398.9
6.0
227.8
-84.6
-22.8
1525.3
24.3
0.4
84.5
276.5
1911.0
o
*Net expansion gas is defined as the energy of isentropic expansion of product gas (stream 7) to 20 psia (1.4 kg/cm^), less the heat required to
maintain the temperature of the expanded gas at the temperature of stream 7 before expansion.
tThis quantity is the net energy transferred and sold to the combined cycle portion of the system.
-------
Table XIV-5
Combined Cycle Case, Sulfur Balance
INPUT
Coal:
6P 1 40 lb wet coal/hr X °'94 'b dry °°al X
Ib wet coal
0.044 IbS molS
Ib dry coal 32.06 Ib S
OUTPUT
Ash:
"t ur it -% -x a0058 lb S x mois
86,446 lb/h,X ^^ X32Q6ibs-
Wastewater to cooling tower
By-product sulfur
Stack gas from sulfur plant (250 ppm as SOj)
Product gas to combined cycle
Total
(gm-mol/sec)
Mol/Hr
0.1
755.7
1.3
17
789.7
(99.5)
%
95.7
0.2
2.1
100.0
SC>2 recycle. The performance of the sulfur removal and re-
covery units is discussed later. The sulfur content of the
treated low-Btu product gas is 184 ppm
sulfur content of the coal.
and is a function of the
The sulfur emitted from combustion of this gas is about
0. 2 lb SC-2/106 Btu (0.36 kg SC>2/106 kcal) of fuel gas. However,
this gas is designed as feed for production of about 1000 MW
(gross) of power when efficient high-temperature, combined
cycle systems are developed. An evaluation of the improved
efficiency of combined cycle power generation is possible by
considering the relative fuel requirements, compared to
In this synopsis the ppm of sulfur in the gas refers to the vol-
umetric concentration of sulfur-containing species, calculated
as monomolecular sulfur content.
XIV-14
-------
conventional steam power plants. Additional fuel would be re-
quired in a conventional plant to attain the power output con-
sidered here. If the emission of the combined cycle facility
was based on the fuel requirement of a conventional plant sized
to the same output, the calculated emissions would be only
0.12 Ib SO2/106 Btu (0. 216 kg SO2/106 kcal) of feedstock.
(7) Gas Purification, Fuel Gas Case
In the fuel gas case, product gas can be supplied at low
pressure for use by industrial users. The Stretford process
was chosen for sulfur recovery because this system will process
the low pressure gas and regenerate elemental sulfur in a single
process.
The 800 degree F (425 degree C) gas from the waste heat
boiler is passed through a high efficiency precipitator to remove
any remaining particulate matter before going to an expansion
turbine. The turbine is operated so that the gas is maintained
above its dew point to avoid condensation within the turbine.
The gas is then cooled from 216 degrees F to 150 degrees F
(102 degrees C to 66 degrees C) before being fed to the Stretford
sulfur recovery unit. This temperature is still above the con-
densation point of the water in the gas; therefore, a separate
wastewater stream is not condensed in this process. Sulfur is
recovered in the Stretford unit (as discussed later) and the sweet
gas (stream 3) is sent to users at a slightly elevated tempera-
ture. Note that the gas cannot be stored or cooled, because an
ammonia-rich liquor will condense.
(8) Energy Balance. Fuel Gas Case
The overall energy balance for the fuel gas case is shown
in Table XIV-6. Conversion of input energy to fuel gas is cal-
culated to be 79. 3 percent. Including by-product power, the ef-
ficiency increases to 82. 5 percent. If the by-product energy
can be delivered as steam, rather than electric power, the over-
all efficiency increases to 87 percent. As discussed in Section (5)
of this chapter, the overall steam, water and electrical require-
ments have not been estimated.
XIV-.15
-------
Table XIV-6
Energy Balance Calculations, Fuel Gas Case
Carrier
Coal
Total energy input
Product fuel gas
By-product power
Sulfur
Ash
Losses
Total energy output
Calculations
612,1401b/hrxl2388Btu/lb
73,988 mol/hr x 378.5 ft3/mol x 214.8 Btu/ft3 (wet basis)
72,570 kWhx 341 3 Btu/kWh
749.7 mol/hr x 32.06 Ib/mol x 3983 Btu/lb
85,81 2 Ib/hrx 3628 Btu/lb
Heat to atmosphere from cooling water, power recovery losses,
plant effluent sensible heat losses
Energy
(x!06Btu/hr)
7583.2
7583.2
6015.4
247.8
95.7
311.3
913.0
7583.2
(x!06kcal/hr)
1911.0
1911.0
1515.9
62.4
24.1
78.5
230.1
1911.0
X
t—I
<
I
o\
-------
(9) Sulfur Balance, Fuel Gas Case
The sulfur balance for this process is shown in Table XIV-7.
The Stretford unit recovers about 95 percent of the sulfur that is
contained in the feed coal as elemental sulfur. The product fuel
gas contains about 3 percent of the initial sulfur in the coal at a
concentration of 325 ppm. Combustion of this gas will emit
0.25 Ib SO2/!06Btu (0.45 kg SO2/lo6kcal) of fuel gas consumed.
This is significantly less than that attainable for direct combus-
tion of the coal, even with stack gas treatment to Federal EPA
New Source Performance Standards.
Table XIV-7
Fuel Gas Case, Sulfur Balance
INPUT
Coal:
n- uoii w-t -,i/h, x°'941bdry coalx
o i z, i tu ID we i coai/ 111 A A
Ib wet coal
0.044 Ib S mol S
Ib dry coal 32.06 Ib S
OUTPUT
Ash:
o. „,„„ , „ „ 0.0059 lbSv mplS
85,812 Ib ash/lir X X
NaCNS produced in Streford tower
Sulfur product
Fuel gas
Total
(gm-mol/sec)
Mol/Hr
789.7
15.7
0.3
749.7
24
789.7
( 99.5)
%
100.0
2.0
94.9
3.0
100.0
XIV-17
-------
3. DISCUSSION OF POLLUTION CONTROL PROCESSES:
COMBINED CYCLE CASE
As discussed in Section 1 of this report, the entire process for
the manufacture of a low-Btu gas may logically be considered as a
pollution control measure since it permits the utilization of high-sulfur
coal in an environmentally satisfactory manner. In the application
proposed here, however, the gas has intrinsic value because it may
be utilized in a combined cycle system for power generation. There-
fore, only those portions of the U-Gas process considered as pollution
control processes are discussed below. The major waste streams are
shown in Table XIV-1, and control treatment processes are summarized
in Table XIV-8.
(1) Control of Coal Dust
Coal dust is generated during the breaking and crushing
operations. This material can be collected with induced draft
systems equipped with cyclones and/or bag filters. The col-
lected particulate matter is recycled to the process.
Table XIV-8
Source and Treatment of Major Waste Streams for U-Gas
Process—Combined Cycle Case (Airblown Gasifier)
Intermediate
Waste Stream
Source
Treatment
1. Coal dust
2. Ash
3. Hydrogen sulfide,
carbonyl sulfide,
hydrogen cyanide,
sulfur dioxide
4. Wastewater
Coal preparation
U-Gas reactors
From gas product, or
removed in acid-gas
treatment and
wastewater treating
Scrubber
Cyclones, bag filters recover
dust for return to process
Agglomerated ash is slurried,
depressurized, dewatered
and returned to mine
Selexol, Claus, and Wellman-
Lord process with an
incinerator
Modified Chevron process,
with ammonia recovery
XIV-18
-------
(2) Ash Collection and Disposal
The ash agglomeration feature of the gasifier produces
pellet-size particles that are removed by slurrying with water.
Pellets drop from the gasifier into a slurry surge pot that is
continuously discharging to lock hoppers for pressure reduction.
After depressuring, the lock hopper discharges by gravity flow
to a screen separator to separate and recycle the water. The
pellets are stored in a surge pile and are eventually discarded
in mined-out areas. The nature of the solids discharged from
the reactor should minimize any leaching from the ash.
(3) Sulfur Recovery
For the combined cycle case, the product gas must be
maintained at pressure. It is also advantageous to leave the
carbon dioxide and water vapor in the gas stream (power re-
covery in the gas turbine is thereby increased). A Selexol acid-
gas removal unit can be designed to remove most of the sulfur
and leave the majority of the carbon dioxide in the product. How-
ever, the process does dehydrate the feed gas.
The Selexol process (developed by Allied Chemical Cor-
poration) is based on the physical solubility of acid-gases in a
glycol ether solvent. The solubility of hydrogen sulfide in this
solvent is much greater than that of carbon dioxide; therefore,
the absorption conditions can be established to remove hydro-
gen sulfide preferentially over carbon dioxide. The resulting
effluent stream (stream 8) containing 16 percent H^S in CO2 is
satisfactory for recovery of sulfur by the Glaus process. The
product gas (stream 7) contains only 50 ppm I^S and the ma-
jority of the original CC>2. One other species of sulfur, car-
bonyl sulfide, must be considered in this system. The solubility
of carbonyl sulfide (COS) in the Selexol solvent (as in most sol-
vents) is intermediate between that of B^S and CO2« Therefore,
an absorber design that preferentially recovers H^S over CO2
cannot recover a high proportion of the COS. In the example
presented, approximately 50 percent of the COS is recovered
with the H2S; according to estimates from licensors of various
solvent-based processes, the remainder stays in the product
gas. The total sulfur content of the treated gas is 184 ppm and,
after combustion, results in an emission of 0. 2 Ib SO2/10^ Btu
(0.36 kg/106 kcal) of heating value in the fuel gas.
XIV-19
-------
The carbonyl sulfide in the treated gas is a major contribu-
tor to the total emitted pollution and deserves further discussion.
In this case, the COS formation within the gasifier was assumed
to follow thermodynamic equilibrium (although the COS concen-
tration may be significantly higher or lower, depending upon the
kinetics of the sulfur reactions). The Selexol process recovers
about half the COS present, so the contribution of COS to final
emissions (on a Btu basis) is directly proportional to the sulfur
content of the coal and independent of the Btu rating of the gas.
Rich H2S off-gas from the Selexol unit is combined with a
sulfur dioxide recycle stream from the Wellman-Lord tail-gas
scrubber and a small gas stream from wastewater treatment.
These sulfur-containing streams are treated in a Glaus unit
that recovers 259 L tons/day (263 m tons/day) of elemental
sulfur. Glaus reactors are discussed in detail in Section 1 of
this report.
Tail-gases from the Glaus plant are incinerated to convert
all sulfur compounds to SO2» so they can be concentrated in the
aqueous sodium sulfite solution of a Wellman-Lord unit. Gas
vented to the atmosphere will contain about 250 ppm SO2; the
recovered SO2 is recycled to the Glaus unit.
The sulfur removal and recovery system described here
may not be ideal. The high-pressure Stretford unit, as dis-
cussed in Chapter XIII could have been employed. It is simpler
and leaves all the CO2 in the product gas but releases greater
emissions. A still more desirable system would treat the gas
at elevated temperature (above 800 degrees F, 425 degrees C)
and leave all the water vapor and carbon dioxide in the product
gas. Several such systems are under development at various
laboratories. Processes for high-temperature desulfurization
are now being developed. One of these will be considered in
Phase II of this study. Although theory dictates that higher
temperature sulfur removal cannot be complete, this potential
process promises much higher efficiency for power generation
from coal.
XIV-20
-------
(4) Wastewater Treatment
The major pollutants in the composite wastewater stream
(stream 3) in the combined cycle case are small amounts of hydro-
gen cyanide, ammonia, carbon dioxide, and hydrogen sulfide. The
composition is reported in Table XIV-1. Hydrogen sulfide in the
wastewater stream is only a small fraction of the total generated
in the process.
The potential pollutants are recovered by a modified
Chevron wastewater treatment. Hydrogen sulfide, hydrogen
cyanide, and some carbon dioxide are removed by a stripping
tower, and sent to the sulfur recovery unit (stream 5). Ammonia
is recovered by a second stripping tower and is dehydrated, per-
haps in a Phosam process, which produces an anhydrous ammonia
by-product. The HCN which is recovered is incinerated in the
sulfur burner of the Glaus plant.
The wastewater, after purification by the modified Chevron
process (stream 4), may still contain trace amounts of ammonia,
cyanide, and sulfur compounds, but the water quality is satisfactory
for cooling water makeup. The low-concent ration contaminants in
the cooling water are eventually eliminated by biological action
in the cooling system, or they are discharged in the cooling
tower blowdown. The latter can be used in such applications as
quenching of hot ashes discharged from the bottom of the gasifier.
Blowdown from cooling towers and boilers, as well as
sanitary sewage and runoff water, can be treated by a miscel-
laneous treatment section incorporating a sewage treatment
plant. The only waste product discharged from the wastewater
treatment stage will be an inert, solid residue.
4. DISCUSSION OF POLLUTION CONTROL PROCESSES;
FUEL GAS CASE
In the fuel gas case, the coal is gasified to generate a low-Btu
fuel gas, perhaps for industrial use. As such, the overall process may
be considered as pollution control because it permits the utilization
of high-sulfur coal in an environmentally satisfactory manner. How-
ever, within the constraints of this project, only the unit pollution
control processes are discussed below. The major waste streams
are shown in Table XIV-9.
XIV-21
-------
The ash and coal dust are handled in the same manner as in the
combined cycle case and are discussed in a previous section of this
chapter. The nitrogen stream from the oxygen plant requires no
treatment. Also, this process is specifically designed to avoid liquid
condensation — another potential pollutant.
Table XIV-9
Source and Treatment of Major Waste Streams for U-Gas
Process — Fuel Gas Case (Oxygen-Blown Gasifier)
Intermediate
Waste Stream
1. Coal dust
2. Ash
3. Hydrogen sulfide,
carbonyl sulfide,
hydrogen cyanide
4. Nitrogen
Source
Coal preparation
U-Gas reactors
f
From gas production
Oxygen plant
Treatment
Cyclones, bag filters return
to process
Agglomerated ash is slurried,
depressurized, dewatered,
and returned to mine
Stretford process
No treating necessary
In this example, the low-Btu product is supplied at low pressure
for use as fuel. A Stretford unit was selected for sulfur removal in
this application. The Stretford process uses a chemical-type solvent
consisting of an aqueous solution containing sodium carbonate, sodium
vanadate, and anthraquinone disulfonic acid (ADA). The feed gas is
countercurrently scrubbed with this solution. The solution reaches
an equilibrium with respect to the CC^; only relatively small amounts
are removed by the process. The same effect is assumed for the
ammonia in the feed gas. Ammonia is shown in the product fuel gas
(stream 3) at 100 ppm. The E^S dissolves in the alkaline solution and
is removed to less than 10 ppmv. The solution is regenerated by air-
blowing and elemental sulfur is removed by froth flotation at a rate of
258 L tons/day (262 m tons/day).
XIV-22
-------
The HCN in the raw gas reacts with sulfur to form sodium thio-
cyanate which may build up and require purging. A modification to the
process is under development to effect HCN removal via reaction with
sodium polysulfide to produce sodium thiosulfate. Additional thiosul-
fate may be generated due to the elevated operating temperature,
thereby requiring treatment of a small purge stream.
The Stretford process is specific for removal of HoS and does
not remove COS that is present in the raw fuel gas. Hence, as in the
combined cycle case, the primary sulfur species contributing to emis-
sion is carbonyl sulfide, and the COS in the product gas is directly re-
lated to the sulfur content of the coal. The emission, on a Btu basis,
will be independent of the fuel heating value, but the sulfur concentra-
tion, on a ppm basis, will decrease as the fuel is diluted with inert
gases. Combustion of this fuel gas will discharge sulfur emissions of
0.25 Ib SO2/106 Btu (0.45 kg SO2/106 kcal). These emissions are sig-
nificantly below those specified by the EPA New Source Performance
Standards of 1.2 Ib SO2/106 Btu (2.16 kg SO2/106 kcal) for the alter-
native, direct combustion of coal.
5. INCREMENTAL COSTS OF POLLUTION CONTROL PROCESSES*
Costs of the pollution control and treatment processes described
above are calculated and presented in this section. The major assump-
tions and conventions adopted are discussed in Chapter III of this report.
Incremental** capital and operating costs are estimated for both
cases and shown in Tables XIV-10, 11, 12 and 13.
The derivation of the formulae for calculating the incremental
cost of gas production due to pollution control is presented in Chap-
ter III of this report. For the discounted cash flow (DCF) method,
The costs presented here are preliminary estimates; the actual
costs of emission control in this process will not be known until
a full-scale facility has been constructed and operated.
Incremental costs in this report refer to that portion of the total
plant cost that is assignable to pollution control facilities.
XIV-23
-------
Table XIV-10
Incremental Capital Costa for Pollution Abatement
for Combined-Cycle Facility Using Airblown
U-Gas Process
Item
$xlO°
Waste heat boiler - 2
Air cooler
Scrubber cooler
Selexol
Claus
Wellman-Lord
Wastewater treating
Particulate emission control
Subtotal installed equipment cost3
Project contingency (15%)
Incremental plant investment (1)
Startup costs (S)c
Interest during construction"
Working capital (W)e
Incremental capital investment (C)
6.6
1.7
0.7
13.0
2.3
4.0
. 3.2
2.2
33.7
5.1
38.7
1.1
DCF Method Utility Method
9.2
1.0
50.0
6.5
0.8
47.1
Notes
Incremental plant investment, return on investment during construction, and
working capital are treated as capital costs in year 0 (the year ending with
completion of startup operation).
"Installed costs, including engineering design costs, contractors' profit and
overhead, and the contingency includes costs for unexpected site preparation
and hardware requirements at 15% of plant investment.
cAt 20% of incremental gross operating cost.
^For the DCF method, computed as the discount rate x incremental plant
investment for 1.875 years' average construction period.
I(l+i)n = I(1.12)1:875= 1.2367611 or 0.23676111 additional investment.
For the utility financing method, computed as the interest rate on debt
x incremental plant investment x 1.875 years.
eSum of materials and supplies at 0.9% of incremental plant investment and
net receivables at 1 /24 of annual incremental revenue.
XIV-24
-------
Table XIV-11
Incremental Operating Costs for Pollution Abatement
for Combined-Cycle Facility Using
Airblown U-Gas Process
Item
Labor
Direct operating labor (DOL)
5 men/shift X $5/hr X 8304 shift hr/man yr
Maintenance labor (1.5% of 1)
Supervision ( 1 5% of direct operating and main-
tenance labor)
Subtotal labor cost
Administration and general overhead
(60% of labor subtotal)
Other direct cost3
Supplies
Operating (30% of DOL)
Maintenance (1.5% of I)
Subtotal supplies
Local taxes and insurance (2.7% of I)
Incremental gross operating costs
By-product credits'5
Sulfur at S10/LT
Ammonia at $25/ton
Total by-product credit
Net operating costs (N)
$/Year
207,600
580,500
118,200
906,300
543,800
2,528,600
62,300
580,500
642,800
1,044,900
5,666,400
851,000
15,000
866,000
4,800,400
Notes:
t materials include:
Purchased utilities:
Makeup boiler feed water at 30
-------
Table XIV-12
Incremental Capital Costa for Pollution Abatement
for Fuel Gas Plant Based on U-Gas Process
(Oxygen-Blown Gasifier)
Item
Incremental sulfur recovery - Stretford
Waste heat boiler to generate steam for Stretford
Exchanger
Modification to Stretford to handle NaCNS
Particulate emission control
Subtotal installed equipment cost
Project contingency ( 1 5%)
Incremental plant investment (1)
Startup costs (S)c
Interest during construction"
Working capital (W)e
Incremental capital investment (C)
SxlO6
12.6
0.2
0.7
1.0
2.2
16.7
1.7
18.4
0.7
DCF
4.4
0.5
24.0
Utility Method
3.1
0.4
22.6
Notes
Incremental plant investment, return on investment during construction, and working
capital are treated as capital costs in year 0 (the year ending with completion of startup
operation).
"Installed costs, including engineering design costs, contractors' profit, and overhead, and
the contingency includes costs for unexpected site preparation and hardware requirements
at 15% of plant investment.
cAt 20% of incremental gross operating cost.
"For the DCF method, computed as the discount rate x incremental plant investment for
1.875 years' average construction period.
I(l+i)n = I(1.12)1:875 = 1.2367611 or 0.2367611 additional investment.
For the utility financing method, computed as the interest rate on debt x incremental
plant investment X 1.875 yrs.
eSum of materials and supplied .9% of incremental plant investment and net receivables
at 1/24 of annual incremental revenue.
XIV-26
-------
Table XIV-13
Incremental Operating Costs for Pollution Abatement
for Fuel Gas Plant Based on U-Gas Process
(Oxygen-Blown Gasifier)
Item
$/Year
Labor
Direct operating labor
(2 men/shift X $5 hr X 8304 shift hr/man yr)
Maintenance labor (1.5% of I)
Supervision (15% of direct operating and maintenance
labor)
Subtotal labor cost
Administration and general overhead
(60% of labor subtotal)
Other direct costs3
Supplies
Operating (30% of Direct operating labor)
Maintenance (1.5% of I)
Subtotal supplies
Local taxes and insurance (2.7% of I)
Incremental gross operating costs
By-product credit (sulfur at $10/LT)b
Net operating costs (N)
83,000
266,000
53,900
402,900
241,700
2,242,600
24,900
266,000
290,900
496,800
3,674,900
- 845,900
2,829,000
Notes:
aOther direct costs:
Steam at $l/10001bs
Power at 1.5
-------
the required incremental annual cost of gas, X, for the assumed rate
of return is:
where
N
I
S
W
X = N + 0.238161 + 0.1275S + 0, 230777W
incremental net operating cost
incremental plant investment
startup costs
incremental working capital
For the utility financing case, it is:
X = N + 0.1198C + 0. 0198W
where
C = required incremental capital investment
N = as defined above
W = as defined above
Assuming an annual gas production of G, the incremental cost of gas,
due to pollution controls is:
annual cost of gas _ X
annual gas production G
Solutions to the above equations for the incremental cost of gas are
shown in Table XIV-14.
Table XIV-14
Incremental Cost of Gas ($/106 Btu)
Accounting Method
DCF
Utility
Combined Cycle Case
Incremental Annual
Cost of Energy
(S/year)
14,371,800
10,458,400
Incremental Cost of
Energy ($/ Ifl6 Btu)
0.33
0.23
Fuel Gas Case
Incremental Annual
Cost of Gas ($/yr)
7,415,800
5,532,400
Incremental Cost
ofGas($/106Btu)
0.16
0.12
XIV- 28
-------
(1) Combined Cycle Case
The capital investment costs of pollution control equip-
ment used in the combined cycle case (airblown gasifier ) are
about $50 million for a system producing 130 x 10^ Btu/day.
Based on the amount of product gas produced, the cost is be-
tween $0. 23-$0.33/IO6 Btu when operating costs are included.
The total costs for this gasification system have also been
estimated for both low temperature desulfurization as well as
for high-temperature desulfurization by the IGT Meissner pro-
cess when it is developed. For low temperature desulfurization,
the overall plant investment is expected to be about $110 million;
pollution control is therefore about 45 percent of the total cost.
Similarly, the incremental increase in gas price resulting from
pollution control is about 50 percent of the total processing cost.
When high temperature desulfurization is developed, the
expected overall price of low-Btu gas is expected to decrease by
about 25 c7lO^ Btu, both because of the simplified cleanup system
and because of the greater thermal efficiency of high-temperature
processing.
(2) Fuel Gas Case
The costs of pollution control in the oxygen-blown fuel gas
case are somewhat lower than for the combined cycle case, pri-
marily because of simplified treatment and elimination of waste-
water treatment. In this case, the incremental capital require-
ment is about $23-24 million and the incremental increase in the
price of gas is about $0.12-$0.16/million Btu, when operating
costs are included.
XIV-29
-------
6. REFERENCES
(1) Loeding, J. W. and Tsaros, C. L., "IGT U-Gas (Clean
Utility Gas) Process, " Clean Fuels from Coal Symposium,
Institute of Gas Technology; Chicago, Illinois (September
10-14, 1973).
(2) Jequier, L. et al., "Apparatus for Dense - Phase
Fluidization, " U. S. Patent 2, 906, 608 (September 29, 1959).
(3) Jequier, L., Longchambon, L., and G. Van De Putte,
"The Gasification of Coal Fines, " Institute of Fuel Journal,
33 (1961), 584-591.
XIV-30
-------
XV. THE KOPPERS-TOTZEK PROCESS
The Koppers-Totzek (K-T) process, developed by the Heinrich
Koppers Gm H of Essen, is a commercially proven scheme to pro-
duce a gas having a heating value of about 2670 kcal/m^ (300 Btu/ft^).*
This gas can be used directly as utility fuel, for chemical synthesis
or as a base for SNG.
In the K-T gasifier, pulverized coal^ reacts with steam and
oxygen (at high temperature and low pressure). Mixing valves blend
the coal, steam and oxygen together prior to entering the gasifier.
All fines can be used as feed and no pretreatment is required. The
gasifier is refractory lined and equipped with a steam jacket to pro-
duce low-pressure steam (reference 3).
Since its first commercial installation in 1952 (in Finland),
16 commercial plants based on the K-T process have been built or
are currently under construction (reference 4).
1. PROCESS DISPLAY
A schematic flow diagram for the K-T gasification process is
shown in Figure XV-1. Stream compositions and flow rates are
given in Tables XV-1 through XV-3 for typical Illinois, Eastern, and
Western coals, respectively. Compositions and heating values for
these coals are shown in Table XV-4. Overall energy blances are
reported in Figure XV-1 and in Table XV-5.
_M—~«^•_««••» O
* In this synopsis, ft refers to one cubic foot when measured at
standard conditions of 60 degrees F, 30" Hg and saturated with
water.
t The K-T gasifier can also be adapted to accept liquid fuels.
XV-1
-------
FIGURE XV-1
The Koppers-Totzek Process
to
BFW - BOILER FEED WATER
CW-COOLING WATER
mton - METRIC TONS
ENERGY BALANCE u 106 kcal'lu)
CARRIER
COAL INPUT
PRODUCT GAS
STEAM
SULFUR
EXHAUST GAS
UNGASIFIED CARBON
DRYING COAL
LOSSES
TOTAL OUTPUT
ILLINOIS COAL
2051
1432
0
34
3
79
22
481
2051
EASTERN COAL
1970
1470
17
6
1
76
19
381
1970
WESTERN COAL
2008
1470
32
5
1
81
25
394
2008
\
•SULFUR CONTENT WHEN OFtlEO TO 2"4 MOSITUHE IN EACH CASE.
-------
Table XV-1
Stream Compositions, Illinois Coal
Stream Names
Stream No.
(Vol.
CO
CO^
M2~
NT
HiS -*- COS
HC1
HO
A
MDEA
Total ( Ib hr I
Dried
Coal
Feed
1
Wt. -:
4.35
0.9-
:.oo
19.08
61.83
4.87
0.179
r>.7:
_
31.08:
6.931
_
I4.:9i
136.333
441.798
34.797
1.279
48.017
_
714.5:8
324.110
Steam
:
r:.338
78.173
Oxygen
)
2.0
08.0
9.956
487.830
497.786
::5.796
Slag
4
68.167
68.16^
30.9: 1
H,O
From
Cooling
Tower
S
1.5:1 :> ""
690.040
Gas
Out
Gasifier
6
51.14
5.60
31.61
0.98
1.77
0.06
8.84
880.950
151.580
38.867
16.887
38.740
1.315
97.887
68.167
::.090
1.316.483
597,157
BFW
to
Gasifier
Jacket
7
153.545
153,545
69,648
L.P.
Steam
from
Jacket
8
146.:34
146.:34
66.33:
Spray
Coulins
Water
9
87.904
87.904
39.873
Gas
to
Waste
Boiler
10
47.38
5. IS
2i) 28
.91
1.64
.05
15.54
880.950
151.580
38.867
16.887
38.740
1.315.
185.79:
68.167
::.09o
1,404,388
637,030
Gas
Out
Waste
Boiler
11
47.38
5.18
:9.:s
.91
1.64
.05
15.54
880.950
151.580
38,867
16.887
38.740
1.315
185.79:
68.167
::.090
1,404,388
637,030
BFW
to
Boiler
12
99:.964
99:.964
450.408
Steam
from
Boiler
13
945.688
945.688
4:8.964
Slag
Cool
H2O
Return
14
1.5:1.250
690.039
Gas
Out
Absorber
15
61.05
_
37.73
1.17
HiS .0:
COS .01
_
-
880.950
—
38.867
16.887
HiS 466
COS 510
_
_
_
_
937.680
4:5.33:
Liquid
Out
Absorber
16
153.8:4
H-iS35 9~18
COS 510
1.085.565
l.:75.827
578.715
Condensate
Out
Cooler
17
1.315
185.792
68.167
22.090
277.364
Gas
Out
Gaus
18
154.198
59.175
19.020
109
105.463
Sulfur
Out
Claus
19
34.032
34.032
15.437
MDEA
10
Absorber
20
1.085.565
1.085.565
492.412
X
-------
Table XV-2
Stream Compositions, Eastern Coal
Stream Names
Sin-am No.
<:.,.(•..„„•„.»„„:
CO-
II;"
\ •
ii-s « cos
H-O
Ash
C
S
°:
Flow rate lib lin
CO
CO-
ll-"
N-
II;S+COS
H-O
Ash
C
S
O-
Slcam
SO-
MD'EA
Total lib hr)
(kc/hn
Dried
Coal
Feed
1
Ul.
4 /'(I
1.37
:.ou
13.7:
• -O.SS
i.os
".05
30.ro
S.456
12.315
64.47.S
430.:"!
6.1.50
43.408
615,7:8
2-9,294
Sleam
2
180.711
180.71 1
86.053
0\yeen
3
20
9S.O
0.560
468.427
477.087
216.815
Slag
4
42.259
42, 239
10.160
H-,0
From
Cooling
Tower
S
956.434
056.434
433.838
Gas
Out
Gasifier
6
50.2"
5.5.S
55.5:
1.05
0.34
0.45
858.000
140.645
40."41
I-. .)«(,
7.41 1
103.486
42.239
21.515
1.241.040
562.936
BFW
to
Gasifier
Jackel
7
I32.0(>2
132.06:
50.003
L.P.
Sleam
from
Jackel
8
i:5.774
i:5.774
57.051
Spray
Cooling
Water
9
84.067
84.967
38.541
Gas
to
Waste
Boiler
10
5.1"
50.-W
O.'K
0.516
15.03
858.009
140.645
40. -41
1 7.996
7.411
188.453
42.230
21.513
1.326.007
601,477
Gas
Out
Waste
Boiler
II
46.62
5.17
30.0-1
O.OS
0.516
1 5.03
858.000
140.645
40.741
17.096
7.557
188.453
42.239
21.513
1,326,153
601,543
BFW
to
Boiler
12
965.260
965.260
437.842
Sleam
from
Boiler
13
910.302
919.302
416.995
Slag
Cool
HO
Return
14
056.436
056.436
433.830
Gas
Out
Absorber
15
2..-1.-
3S."1
1.20
H-S .021.
COS .003
858.009
66. (.(.0
40.741
17.006
H-S 4(.v*
COS UiO
983,975
446,331
Liquid
Out
Absorber
16
85.000
II -5 6.484
COS 100
90.184
40.907
Condensate
Out
Cooler
17
188.455
42.239
21.513
25:.:o.-
114,400
Gas
Out
Gaus
18
S3.600
10.644
3.421
40
97,705
44,310
Sulfur
Out
Claus
19
6.126
6.126
2. "70
MDEA
to
Absorber
20
498 000
498.000
225.SO!
X
-------
Table XV-3
Stream Compositions, Western Coal
Stream Name
Stream No.
Gas composition
( VolT 1
CO
CO,
H,~
N;
H;S + cos
HC1
H-,0
Alh
C
s
Ch
o-T
Steam
Flow rate (Ib/hr)
CO
CO,
H-,
N;
H;S + cos
~
HCI
H,O
Ash
C
S
Cl,
°2
Steam
SO,
MOEA
Total (Ib/hr)
(kg/hr)
Dried
Coal
Feed
I
wCZ
4.24
1.01
-
2.00
2:. 14
56.74
0.67
0.035
13.18
_
-
34.163
8.138
_
16.115
178.394
457.182
5.398
282
106.197
805,869
365.542
Steam
2
84.894
84.894
38.508
Oxygen
3
2.0
98.0
10.441
511.607
522,048
236,801
Slag
4
89,197
89.197
40.460
H,0
From
Cooling
Tower
S
1. 990.559
1,990.559
902.918
Gas
Out
Gasifier
6
55.07
5.68
30.03
1.11
0.28
0.01
7.80
918,372
148.852
35,769
18,579
6,044
290
83.653
89,197
22.859
1,323,615
600,392
BFW
To
Jacket
7
173.047
173.047
78.494
Steam
From
Jacket
8
164.807
164.807
74,756
Spray
Cooling
Water
9
85.470
85,470
38.769
Gas
To
Boiler
10
51.00
5.26
27.81
1.03
0.26
0.01
14.61
918.372
148,852
35.769
18.579
6.044
290
169.123
89.197
22.859
1,409,085
639,161
Gas
Out
Boiler
11
51.00
5.26
27.81
1.03
0.26
0.01
14.61
918.372
148,852
35.769
18.579
6.044
290
169.123
89,197
22,859
1,409,085
639,161
BFW
To
Boiler
12
968,976
968,976
439,528
Steam
From
Boiler
13
922,843
922.843
418,602
Slag
Cooling
H,O
Return
14
_
1,990,558
1,990,558
902,917
Gas
. Out
Absorber
IS
61.61
3.51
33.60
1.24
H-,5 .025
COS .0027
_
_
918,372
82,192
35,769
17,859
H,S454
C6S89
-
—
1,054,735
478,428
Liquid
Out
Absorber
16
67.056
H,S5I81
COS 89
399,336
471,662
213,946
Conden-
sate
Out
Cooler
17
290
169,123
89.197
22,859
281,469
127,674 '
Gas
Out
Claus
18
67,121
8,534
2.743
32
78,430
35,576
Sulfur
Out
Glaus
19
4,908
4,908
2,226
MDEA
To
Absorber
20
399,336
399,336
181,139
-------
(1) Basis for Analysis
The process flow sheet data, stream composition and
flow rate data presented are based on a facility scaled to pro-
duce sufficient product gas to supply 140 x 10^ Btu/day (35, 280
x 10 kcal/day). The basic data underlying this analysis (utiliz-
ing Western, Illinois and Eastern coals) was supplied by the
Koppers Company. Each of the coals is first crushed to 70 per>
cent less than 200 mesh, and then dried to 2 percent moisture
(by weight) for use as feed to the gasifier.
(2) Layout and Symbols
The general direction of process flow is from the coal
supply on the left to the product gas on the extreme right of the
flow design with residuals and by-products shown along the
bottom. The bold line indicates the flow of the primary gasifi-
cation process. The encircled figures refer to the stream coin-
positions shown in Tables XV-1, XV-2, and XV-3.
Table XV-4
Coal Composition and Heating Values
Carbon
Hydrogen
Nitrogen
Sulfur
Oxygen
Chlorine
Water (moisture)
Ash
Total
Heating value,
x 106Btu/hr
(x 106kcal/hr)
Western Coal
(Wt %)
56.74
4.24
1.01
0.67
13.18
0.035
2.00
22.14
100.00
7968
2008
Illinois Coal
(Wt %)
61.83
4.35
0.97
4.87
6.72
0.179
2.00
19.08
100.00
8137
2051
Eastern Coal
(Wt %)
69.88
4.90
1.37
1.08
7.05
0.0
2.00
13.72
100.00
7817
.1970
XV-6
-------
Rhombic-shaped units represent intermediate products
(such as boiler feed water), uses (such as electric power), and
sources (such as process return water), for which the distribu-
tion is not known. Nonintegral pollutant cleanup processes are
indicated by sloping rectangles. The extensive treatment and
recovery processes which this symbol represents are discussed
in the process description and pollution control sections of this
process summary. Inverted trapezoids denote residual storage.
2. PROCESS DESCRIPTION
The unit processes which comprise the K-T coal gasification
process to produce low-Btu fuel gas are described in this section.
The generation and disposition of pollutant streams are discussed
when appropriate. The discussion of pollution control processes is
covered in the next section.
(1) Coal Preparation
Coal feed for use in the K-T process can contain 2 to 8
percent moisture (by weight) and must be pulverized to about
70 percent less than 200 mesh (<74/im) in the coal preparation
system. All fines can be fed to the gasifier. For purposes of
this report, each of the three coals is assumed to be dried to
2 percent moisture (by weight) for gasification. The energy
requirement for this drying has been charged to the system in
each case.
(2) Gasification
Pulverized coal is conveyed to gasifier feed bins and
then through screw feeders to the gasifier. Pulverized coal,
oxygen, and low-pressure steam are fed through mixing noz-
zles where moderate temperature and high burner velocity
prevent reaction until entry into the gasifier. The high gasifi-
cation temperature of 3300 degrees F (1815 degrees C) com-
pletely gasifies the carbon and sulfur in the feed and produces
a clean ash in the form of a molten slag. About one-half of
this slag leaves the gasifier with the gas while the other half
XV-7
-------
drops into a water quench tank. Excess heat from the gasifier
reactions is removed by production of low-pressure steam in
the gasifier jacket (stream 8).
(3) Gas Quench and Cooling
Gas removed from the gasifier (stream 6) is quenched with
a water spray (stream 9) to solidify entrained slag droplets. It
is then passed through a waste heat boiler where superheated
steam (stream 13) is produced. The quenched gas (stream 11)
is cleaned and further cooled in a two-stage venturi scrubbing
system (gas cooler). Final particulate removal can be accom-
plished by electrostatic precipitators (not shown).
(4) Gas Purification
The type of gas cleaning system employed in the process
depends primarily on the end use of the product gas. For low
pressure utility or synthesis gas production, a chemical reaction
system (reference 5) using a tertiary amine (methyl-diethanol-
amine-MDEA) will produce a clean fuel gas of relatively low
sulfur content. The amine, at low temperature and pressure,
absorbs essentially all the hydrogen sulfide leaving only 250
ppmv in the gas. According to data provided by the Koppers
Company, the carbonyl sulfide (COS) is both physically ab-
sorbed and chemically reacted in this purification step. About
half of the COS is hydrolyzed to H2S and half of the remainder
is removed by the amines. Therefore, approximately 25 per-
cent of the original COS remains in the treated gas. The MDEA
amine may be blended with other amines to optimize ratios of
CO2, H2S, and COS. For purposes of this study, the CO2/H2S
ratio fed to the Glaus plant (sulfur recovery) is assumed to
be 10; during gasification of high sulfur Illinois coal, even
though essentially all the CO2 is absorbed, a more concentrated
H2S stream is sent to the Glaus plant.
A Stretford process is another viable alternative that
could have application here for sulfur removal. If the K-T
off-gas was compressed for use in a high pressure system
(e.g., to generate SNG), physical absorption acid-gas removal
XV-8
-------
unit processes (such as Rectisol, or a hot potassium carbonate
process) would have to be considered.
(5) Water Supply and Waste Heat Recovery
Makeup water to the system can be supplied by wells, sur-
face streams, or reservoirs. Conventional water treating tech-
niques are used to make high purity boiler feed waters and other
makeup waters for plant needs. The raw water requirements
for the 140 x 109 Btu/day facility vary from about 17, 050 liters/
min (4500 gpm) for the Eastern coal to about 25, 350 liters/min
(6700 gpm) for a Western coal.
Particulate-laden waters from the gas cleaning and cooling
system are sent to a clarifier (streams 14 and 17) where the
sludge is either filtered or sent to the plant disposal area.
Clarified water is reused in the venturi scrubbing system and
in the cooling tower system while the solid residue is returned
to the mine for disposal along with the slag. Evaporation,
windage, and blowdown of the cooling tower, and water in the
clarifier sludge and recovered slag are the major losses of
water in the K-T process. In arid areas where water supply
is at a premium, air cooling would be used wherever feasible.
(6) Oxygen Supply
The (98 percent purity) oxygen required to fire the K-T
gasifier for production of utility gas ranges from 5113 m tons/day
(5636 tons/day) for Eastern coal to 8772 m tons/day (9669 tons/
day) for Western coal. The oxygen is manufactured onsite
using commercially available air separation equipment and
process steam. Water condensed during compressing and
cooling the air is used in the process. The separated nitrogen
can also be used in the plant wherever an inert gas is required
(e.g., transport medium for prepared coal); however, it is
anticipated that the bulk of this stream will simply be returned
to the atmosphere.
The energy required for the oxygen plant ranges from
433, 650 kg/hr (956, 000 Ib/hr) of steam for the Eastern coal to
473, 600 kg/hr (1, 044, 000 Ib/hr) of steam for the Western coal.
XV-9
-------
This steam is used for compression energy within the oxygen
plant, rather than heat energy.
(7) Steam and Power Generation
Temperature control in the K-T gasifier is accomplished
by removal of excess heat. This heat raises low-pressure
steam (stream 8) while superheated steam (stream 13) is ob-
tained by cooling effluent from the gasifier in a waste heat
boiler. The low-pressure steam provides the steam required
for gasification. This by-product steam production can supply
the total steam requirement of the facility (about 36, 000 •* 86, 000
kg/hr); extra boiler capacity is usually not required. Addition-
ally, the oxygen plant requirement is 435, 000 -» 474, 000 kg/hr.
Even including other steam loads, in this analysis, by-product
steam production was found to be in excess of process needs.
For the Eastern and Western coals, 308 and 320 x 106 kcal/hr
(1221 and 1269 x 10° Btu/hr) of steam was produced respec-
tively; process requirements were only 293 and 288 x 10° kcal/
hr (1161 and 1143 x 106 Btu/hr). In the case of the Illinois coal,
however, the process was steam-deficient, producing 327 x 10"
kcal/hr (1298 x 106 Btu/hr) while requiring 360 x 106 kcal/hr
(1432 x 10^ Btu/hr). The deficiency was caused by the much
higher steam requirements of the acid-gas removal system in
this case. Product gas was charged (financially) for the pro-
cess, when using the Illinois coal, to account for this steam
deficiency. However, a more satisfactory system might be
developed by modifying the acid-gas treatment system. A
boiler house and its resulting stack-gas emissions were there-
fore not included in this analysis.
Because these facilities do not require a separate boiler
house, electric power was assumed to be purchased. Also,
power requirements were not supplied by the process licensor
for each of these cases. If we assume that 25 MW of electrical
power is required, the resulting power plant for onsite electri-
cal power generation would be too small to be economical.
Therefore, because the power house for electrical requirements
would be small and because no additional steam is needed in the
process, no power house was included and electrical power was
assumed to be purchased.
XV-10
-------
(8) Energy Balance
Energy balances for gasification of the three coals ana-
lyzed for feeds in a K-T process are shown in Table XV-5.
From this data, the overall process efficiencies, coal-to-gas,
are calculated to be 69.8 percent, 74. 6 percent and 73.2 per-
cent for Illinois, Eastern, and Western coals, respectively.
The lower efficiency for the Illinois coal is caused by the higher
steam demand as discussed in the preceding section. These
energy balances do not include the electrical input to the system
as discussed in the footnote of Table XV-5.
(9) Sulfur Balance
There are essentially only three sulfur-containing effluent
streams generated in the Koppers-Totzek process: the product
gas (stream 15), sulfur recovered from the Glaus plant
(stream 19), and the exhaust gases from the Glaus tail-gas
cleanup system (stream 18). The relative quantities of sulfur
in each of these streams for the Illinois, Eastern, and Western
coals are listed in Table XV-6. Based on expected performance
of the MDEA scrubbing system, the level of E^S in the product
gas was assumed at 250 ppmv.* COS will contribute additional
sulfur to the stream as indicated in Table XV-6. The Glaus
plant off-gas will also contain about 250 ppm sulfur (as 802).
These fixed factors force a higher apparent efficiency for the
Glaus plant when used with Illinois coal. This improved effi-
ciency results because SC>2 is recycled to the Glaus plant from
the tail-gas cleanup system (see Figure XV-1). Higher sulfur
emissions are to be expected with Illinois coal because of the
higher initial sulfur level in the coal (resulting in higher overall
levels of COS in the product gas and greater mass of SO2 in the
cleaned Glaus tail-gas); however, the fraction of sulfur recovery
is greater in this case.
* In this synopsis the ppm of sulfur in the gas refers to the
volumetric concentration of sulfur-containing species, calcu-
lated as monomolecular sulfur content.
XV-H
-------
Table XV-5
Energy Balances for Various Coal Feeds*
Carrier
Coal feed
Total energy input
Product eas
Less product gas for plant fuelT
Heat in excess steam77
Sulfur
Sensible heat in sulfur plant exhaust gas
Heat in unsasified carbon
Heat in molten slasi
Heat for (Irving coal
Losses
Total energy output
Process thermal efficiency
Illinois
x 106Btu/hr
8137.00
8137.00
5833.33
-153.15
0
136.:0
13.80
311.64
0.42
87.32
1907.44
8137.00
Coal
x 106kcal/hr
2050.52
2050.52
1470.00
-38.59
0
34.32
3.48
78.53
0.11
22.00
480.67
2050.52
5833.33-153.15
8137
Easter
x 106 Btu/hr
7817.00
7817.00
5833.33
0
66.00
24.53
5.33
303.50
0.28
75.25
1508.78
7817.00
5833.33
7817 "'
n Coal
x I06kcal/hr
1969.88
1 969.88
1470.00
0
16.63
6.18
1.34
76.48
0.07
18.96
380.22
1969.88
00 = 74.62
Westen
x 106 Btu/hr
7968.00
7968.00
5833.33
0
126.12
19.65
4.28
322.48
0.53
98.49
1563.12
7968.00
5833.33
7968.00 X
iCoal
x I06kcal/hr
2007.94
2007.94
1470.00
0
31.78
4.95
1.08
81.26
0.13
24.82
383.92
2007.94
00 = 73.25
'The electrical power requirements are not included in this energy balance because they have not been specifically developed for each of these cases.
If the power requirements were 25 MW, generated onsite. the gross plant efficiency will drop by about 2 percent.
fCharged as plant fuel to raise sufficient process steam.
•f+Steam generated in excess of plant requirements.
-------
Of the sulfur contained in the feed for the three coals,
the amount remaining in the product fuel gas ranged between
2 and 8.8 percent while 90.9 to 97.8 percent is recovered in
its elemental form for sale. The Wellman-Lord tail-gas treat-
ment of Glaus plant off-gas will contribute between 7 to 24 kg/hr
of sulfur to the atmosphere. This accounts for 0.15 to 0.30 per-
cent of the total sulfur in the original feedstock.
Table XV-6
Sulfur Balances*
Carrier
Input
Coal (Veil
Total input •
Output
In product gas
As l^S
As COS
As Liquid Sulfur
In tail-gas as SOo
Total output
kg/hr
15.784
15.784
19')
123
15.438 •
24
15.784
Illinois Conl
(Ib/hr)
34.797
34.797
439
272
34.032
5-r
34.797
%
100.00
100.00
1.26
0.78
97.8
0.15
100.00
kg/hr
3.016
3,016
200
24
2.783
9
3.016
Eastern Coa
(Ib/hr)
6.650
6.650
441
53
6.136
20
6.650
1
%
100.00
100.00
6.63
0.80
92.27
0.30
100.00
kg/hr
2,449
2.449
194
21
2.227
7
2.449
Western Cua
(Ib/hr)
5,398
5,398
427
47
4,908
16
5,398
1
%
100.00
100.00
79.1
0.87
90.92
0.30
100.00
*For Koppcrs-ToUck plants producing 35,280 x I0fl kcal/day (140 x 109 Btu/day) oflow-Btu fuel gas.
(10) General Comments
The product gas from the K-T process is composed pri-
marily of carbon dioxide, carbon monoxide, and hydrogen with
only trace amounts of methane. The heating value of the product
gas is 2572 kcal/m3 (289 Btu/ft3) for the Eastern and Western
coals; for the Illinois coal, the heating value is 2658 kcal/m3
(298. 7 Btu/ft3) because more CO2 has been removed in the
MDEA processing. The product gas is assumed to be used di-
rectly as a fuel gas. The acid-gas removal system, however,
can be designed to remove essentially all the CC>2 and H^S, so
that the product gas would also be suitable for methanation to a
SNG product gas.
XV-13
-------
3.
Each K-T gasifier unit can process coal feed rates up to
770 m tons/day (850 short tons/day) and hence produce about
1,250,000 m3/day (45 x 106 ft3/day) of 75 kcal (300-Btu) gas
(reference 1). Thus, approximately 11 gasifiers operating in
parallel would be required to supply the 35,280 x 10° kcal/day
(140 x 109 Btu/day) specified as the basis for this study. While
operating in parallel, the gasifiers could share common utili-
ties, water cleanup system, and gas cleanup system.
DISCUSSION OF POLLUTION CONTROL PROCESSES
The possible pollution sources from the K-T process are limited
by the high gasification temperatures (1815 to 1925 degrees C) to the
slag, the process wastewater, and the product gas. The nature and
treatment of these major waste streams are discusser) in this section
and shown in Table XV-7.
Table XV-7
Nature and Treatment of Major Waste Streams
Intermediate Waste
Stream
Source
Treatment
Coal dust and fines
Slag .
Wastewater
H2S, COS, SO2, sulfur
Coal crushing and drying
Gasification and washing
of product gas
Slag quench, venturi
scrubbers, gas washing
Gasification
Cyclone separators, bag
filters enclosure; fed to
gasifier
Clarifier-settler, sludge to
landfill or mine-fill
Biological
Glaus plant, Wellman-Lord
tail-gas cleanup; sulfur
sold
XV-14
-------
(1) Coal Preparation
Pulverization and drying of the coal is to take place in
totally enclosed equipment so that all dusts in any of the efflu-
ent streams can be separated in cyclones, bag filters, or pre-
cipitators. All these fines can be used as feed for the gasifier.
Coal temperatures will be maintained at 80 degrees C (180 de-
grees F) during coal drying so that no chemical reactions or
devolatilization of the feed will occur. The drying gases are
not contaminated and can be discharged to the atmosphere after
separation from the dried and pulverized coal.
(2)
The coal ash is completely slagged* in the gasifier so
that the effluent streams contain minimal amounts of dusts and
leachable materials. Particulate matter (entrained slag) leav-
ing the gasifier with the product gas (stream 11) is cooled and
collected in venturi scrubbers. This source of solids contains
unburned carbon amounting to approximately 5 percent of that
charged to the gasifier. Solids [quenched slag (stream 4) and
particulates] separated from the process waters can be col-
lected and sent to landfill or mine-fill operations. Since the
particulate matter is also slagged material, there should be no
leachable contaminants and disposal should pose no environ-
mental problems.
(3) Process Waste waters and Gas Purification
Removal of particulate matter from the product gas
occurs in the primary and secondary venturi scrubbers where
the gas is contacted with recirculated cooling water. In the
scrubbers, soluble ammonia, hydrogen cyanide, and some
acid-gases are absorbed into the water. Because of the high
gasification temperatures of 1815 to 1925 degrees C (3300 to
Heated to the molten state.
XV-15
-------
3500 degrees F), phenols, pyridines and organics are not
formed and the quantities of other soluble contaminants are
minimized. A bleed stream from the recirculated water sys-
tem is continuously fed to the stripper where gaseous NHo,
CC>2, H2S, and some cyanide are removed and sent to the Glaus
unit for incineration. In the Glaus unit, I^S is catalytically
converted to elemental sulfur. SC>2 contained in the off-gas is
treated in a Wellman-Lord tail-gas process. The stripped
water is cooled and returned to the recirculation system.
Excess water can be bled to the boiler feedwater system for
treatment and generation of steam at the gasifier.
Table XV-8 lists data (reference 1) collected at a
Kutahya, Turkey, Koppers-Totzek plant in 1972; these data
indicate the orders of magnitude for components of interest
and possible contaminants. Although specific tests were con-
ducted, pollutants such as phenols, tars, oils, and pyridines
were not detected.
(4) Product Gas
Even though the crude product gas is washed in venturi
scrubbers and coolers, most of the hydrogen sulfide and carbonyl
sulfide generated in the gasifier remain in this gas because of
their low solubilities in water. The tertiary amine absorption
system leaves only 2 percent (with Illinois coal) to 9 percent
(with Western coal) of the total feed sulfur as H2S and COS in
the product gas. These results are a consequence of the design
parameter (for this study) of 250 ppmv H2S in the product gas.
Operating conditions with the amine system can be altered to
set H2S concentration in the product gas to other levels, within
the operating range of the system.
The product gases delivered by the process (stream 15) at
the design conditions set in this study are clean-burning fuels.
Flue gases resulting from combustion of the K-T product gases
will contain 0.45 kg SO2/106 kcal (0.25 Ib SO2/10 Btu) of
heating value in the fuel gas with Illinois coal, and 0.31 kg
SO2/106 kcal (0.17 Ib SO2/106 Btu) for both the Eastern and
and Western coals. These emissions compare favorably with
the EPA New Source Performance Standards of 2.16 kg SO2/106
kcal (1.2 Ib SO2/106 Btu for the alternative of direct combustion
of coal with stack gas scrubbers.
XV-16
-------
Table XV-8
Koppers Coal Gasification Water Analyses,
Kutahya, Turkey
Sample Location
pH value
Conductivity
CaO
MgO
Ma
K
Zn
Fe
NH4
NO2
NO3
PO4 total
Cl
S04
CN
H2S
KMnO4 consumed
COD
SiO2
Suspended solids
Hot residue, 800°C
Stripped residue
Hot residue, 800°C
Cu
S
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
mgO2/l
mg/1
mg/1
mg/1
mg/1
mg/1
mg/1
I
8.8
7.6
.001
78
97
17.5
5.6
0.01
0.05
0.32
0.02
58.2
1.89
18
42
0.26
8
14
14.8
14
4
568
268
0.01
II
8.8
1.8
.01
101
161
17.5
8.8
0.03
0.22
157
0.13
3.32
0.81
85
216
0.52
9
18
16.0
4612
3918
812
550
0.01
III
8.9
2.0
.01
78
194
17.5
10.0
0.02
1.95
184
4.47
13.7
1.21
96
155
12.5
Not detei
400
128
14.8
5084
4356
940
588
0.01
IV
8.8
1.8
.01
135
145
17.5
8.0
0.02
0.20
137
0.24
24.7
0.81
57
255
1.4
;ted-
11
16
19.8
3072
2690
706
526
0.01
V
8.9
1.8
.01
179
113
17.5
8.0
0.02
0.64
122
4.37
22.9
2.70
46
109
14.0
__
145
63
42.6
50
46
724
512
0.06
I. Cooling water to slag quench tank (stream 5).
II Water from slag quench tank (stream 14).
Ill Wash water from gas cooler (stream 17).
IV Total water to clarifier.
V Water out of clarifier.
XV-17
-------
4. COSTS OF POLLUTION CONTROL*
Costs of the pollution control and treatment processes de-
scribed above are calculated and presented in this section. The
major assumptions and conventions adopted here are discussed in
Chapter III of this report.
The incremental** capital cost and incremental gas price when
using Illinois coal are presented in Tables XV-9a and XV-9b. Similar
data for the Eastern coal and the Western coal are summarized in
Tables XV-10 and XV-11, respectively.
The total incremental capital requirement of the K-T process
for the lower sulfur content Eastern and Western coals is about
$15 million; the higher sulfur Illinois coal requires about a $25 mil-
lion capital investment for pollution control, primarily because of
larger acid-gas removal and sulfur recovery systems.
The incremental increase in gas price resulting from pollution
control in the K-T process is about $0,10 to $0.15/106 Btu for the
lower sulfur coals; with the Illinois coal, the cost is about $0.20 to
$0.25/10^ Btu, depending upon the accounting approach utilized. The
higher cost in the case of the Illinois coal is caused by the greater
capital requirements and the higher process steam demands (included
in "other direct costs").
The derivation of the formulae for calculating the incremental
increase in the cost of gas production due to pollution control require-
ments is presented in Chapter III of this report. Using the discounted
cash flor (DCF) method, the required incremental annual cost of gas,
X, for the assumed rate of return is:
X = N + 0.238161 + 0.0175S + 0.230777W
The costs presented here are preliminary estimates; the actual
costs of emission control in this process will not be known until
a full-scale facility has been constructed and operated.
Incremental costs in this report refer to that portion of the total
plant cost that is assignable to pollution control facilities.
XV-18
-------
Table XV-9a
Incremental Capital Costa for Pollution Abatement for
Koppers-Totzek Process Using Illinois Coal
Item
Gas desulfurization — MDEA
Clans plant
Wellman-Lord tail-gas cleanup
Particulate emission control
Wastcwater treating
Installed equipment cost
Project contingency (15%)
Incremental plant investment (I)
Startup costs (S)c
Interest during construction
Working capital (W)e
Incremental capital investment (C)
$ 10C
4.2
3.2
4.0
3.6
17.2
2.6
19.8
0.5
DCF
4.7
0.7
25.7
Utility
3.3
0.6
24.4
Notes:
incremental plant investment, return on investment during construction,
and working capital are treated as capital costs in year 0 (the year ending
with completion of startup operation).
"Installed costs, including engineering design costs, contractors' profit,
and overhead, and the contingency includes costs for unexpected site
preparation and hardware requirements at 15% of plant investment.
v cAt 20% of incremental gross operating cost.
"For the DCF method, computed as the discount rate x incremental plant
investment for 1.875 years' average construction period.
1(1 + i)n = K1.12)1:875 = 1.2367611 or 0.2367611 additional investment.
For the utility financing method, computed as the interest rate on debt
x incremental plant investment x 1.875 yrs.
eSum of materials and supplies at .9% of incremental plant investment and
net receivables at '1/24 of annual incremental revenue.
XV-19
-------
Table XV-9b
Incremental Operating Costs for Pollution Abatement
for Koppers-Totzek Process Using Illinois Coal
Item
$
Labor
Direct operating labor (4 men/shift x $5/hr.
x 8304 shift hr/man yr)
Maintenance labor (1.5% of I)
Supervision (15% of direct operating and
maintenance labor)
Subtotal labor cost
Administration and general overhead
(60% of labor subtotal)
Other direct costs*
Supplies
Operating (30% of direct operating labor)
Maintenance (1.5% of I)
Subtotal supplies
Local taxes and insurance (2.7% of I)
Incremental gross operating costs
By-product credit (sulfur at S10/LT)
Net operating cost (N)
166,000
297,000
69,500
99,800
297,000
532,500
319,500
6,059,100
346,800
504,600
1,197,700
7,792,500
-1,197,700
6,594,800
*Other direct costs include catalysts and chemicals expended and utilities purchased for the
pollution control processes. The high charges in this case reflect the substantially increased
cost of the steam requirements for the acid-gas treatment of the Illinois coal.
XV-20
-------
Table XV-lOa
Incremental Capital Costa for Pollution Abatement for
Koppers-Totzek Process Using Eastern Coal
Item
Gas desulfurization - MDHA
Clans plant
Wellman-Lord tail-gas cleanup
Particulatc emission control
Wastewater treating
Installed equipment cost
Project contingency ( 1 5%)
Incremental plant investment (I)'1
Startup costs (S)c
Interest during construction
Working capital (W)e
Incremental capital investment (C)
$ 106
1.9
I.I
1.8
-> -i
3.6
10.6
1.6
12.2
0.7
DCF
2.9
0.4
16.2
Utility
2.1
0.3
15.3
Notes:
alncremental plant investment, return on investment during construction,
and working capital are treated as capital costs in year 0 (the year ending
with completion of startup operation).
"Installed costs, including engineering design costs, contractors' profit,
and overhead, and the contingency includes costs for unexpected site
preparation and hardware requirements at 15% of plant investment.
°At 20% of incremental gross operating cost.
"For the DCF method, computed as the discount rate x incremental
plant investment for 1.875 years' average construction period.
1(1 + i)n = I(1.12)1:875 = 1.2367611 or 0.2367611 additional
investment.
For the utility financing method, computed as the interest rate on
debt x incremental plant investment x 1.875 yrs.
eSum of materials and supplies at .9% of incremental plant investment
and net receivables at 1/24 of annual incremental revenue.
XV-21
-------
Table XV-lOb
Incremental Operating Costs for Pollution Abatement for
Koppers-Totzek Process Using Eastern Coal
Item
$
Labor
Direct operating labor (4 men/shift x $5/hr
x 8304 shift hr/man yr)
Maintenance labor (1.5% of I)
Supervision (15% of direct operating and
maintenance labor)
Subtotal labor cost
Administration and general overhead
(60% of labor subtotal)
Other direct costs*
Supplies
Operating (30% of direct operating labor)
Maintenance (1.5% of I)
Subtotal supplies
Local taxes and insurance (2.7% of I)
Incremental gross operating costs
By-product credit (sulfur at $ 10/LT)
Net operating cost (N)
166,000
183,000
52,400
49,800
183,000
401,400
240,800
2,740,600
232,800
329,400
215,500
3,945,000
- 215,500
3,729,500
*Other direct costs include catalysts and chemicals expended and utilities (e.g., cooling water,
electrical power) purchased.
XV-22
-------
Table XV-lla
Incremental Capital Costa for Pollution Abatement for
Koppers-Totzek Process Using Western Coal
Item
Gas dcsulfurization - MDEA
Claus plant
Wellman-Lord tail-gas cleanup
Particulate omission control
Wastcwater treating
Installed equipment cost
Project contingency ( 1 5%)
Incremental plant investment (I)"
Startup costs (S)c
Interest during construction
Working capital (W)e
Incremental capital investment (C)
$ 106
1.5
1.0
1.9
2.2
3.6
10.2
1.5
11.7
0.7
DCF
2.8
0.4
15.6
Utility
2.0
0.3
14.7
Notes:
alncremental plant investment, return on investment during construction,
and working capital are treated as capital costs in year 0 (the year ending
with completion of startup operation).
"Installed costs, including engineering design costs, contractors' profit,
and overhead, and the contingency includes costs for unexpected site
preparation and hardware requirements at 15% of plant investment.
cAt 20% of incremental gross operating cost.
°For the DCF method, computed as the discount rate x incremental plant
investment for 1.875 years' average construction period.
1(1 +i)n = I(1.12)1:875 = 1.2367611 or 0.236761 1 additional investment.
For the utility financing method, computed as the interest rate on debt
x incremental plant investment x 1.875 yrs.
eSum of materials and supplies at .9% of incremental plant investment
and net receivables at 1/24 of annual incremental revenue.
XV-23
-------
Table XV-lib
Incremental Operating Costs for Pollution Abatement for
Koppers-Totzek Process Using Western Coal
Item
$
Labor
Direct operating labor (4 men/shift x $5/hr
x 8304 shift hr/man yr)
Maintenance labor (1.5% of I)
Supervision (15% of direct operating and
maintenance labor)
Subtotal labor cost
Administration and general overhead
(60% of labor subtotal)
Other direct costs*
Supplies
Operating (30% of direct operating labor)
Maintenance (1.5% of I)
Subtotal supplies
Local taxes and insurance (2.7% of I)
Incremental gross operating costs
By-product credit (sulfur at S10/LT)
Net operating cost (N)
166,000
175,500
51,200
49,800
175,500
392,700
235,600
2,201,100
225,300
315,900
172,800
3,370,600
- 172.800
3,197,800
*Other direct costs include catalysts and chemicals expended and utilities (e.g., cooling water,
electrical power) purchased.
XV-24
-------
where
N = incremental net operating cost
I = incremental plant investment
S = startup costs
W - incremental working capital.
For the utility financing case, it is:
X = N + 0. 1198C + 0.'0198W
where
C = required incremental capital investment
N = as defined above
W = as defined above.
Assuming an annual gas production of G, the incremental cost of gas,
due to pollution controls is:
annual cost of gas _ X
annual gas production G
Solutions of the above equations for the incremental cost of gas using
the three coals analyzed as possible feeds in this process are shown
in Table XV-12.
Table XV-12
Incremental Cost of Gas
Accounting Method
IXT
Utility l-'in:ii]dng
Illinois Coal
Incremental Annual
Cost of Cas($/yrl
1 1.5.15.750
'J.SOS.XOO
Incremental Cost
ofCas($/!06Btu)
0.24
0.1"
Eastern Coal
Incremental Annual
Cost of Gas($/yr)
6.81(1.410
5.568.380
Incremental Cost
ofCas($/106Btu)
0.14
0.11
Western Coal
Incremental Annual
Cost of Cast S/yr)
6.165.900
4,964.800
Incremental Cost
ofGas($106Btu»
0.1.1
0.10
XV-25
-------
5. REFERENCES
(1) Farnsworth, J. F., Kamody, J. F., and Mitsak, P. M.,
"Clean Environment With K-T Process. " USEPA:
Environmental Aspects of Fuel Conversion Technology,
St. Louis, Missouri, May 1974.
(2) "Evaluation of Pollution Control in Fossil Fuel Conversion
Processes," Gasification; Section 1: Koppers-Totzek
Process, EPA-650/2-74-009a, January 1974.
(3) Farnsworth, J. F., Leonard, H. F., Mitsak, D. M., and
Wintrell, R., "Production of Gas from Coal by the Koppers-
Totzek Process. " IGT Clean Fuels from Coal Symposium,
Chicago, Illinois, September 1973.
(4) Farnsworth, J. F., Leonard, H. F., Mitsak, D. M., and
Wintrell, R., "The Production of Gas from Coal Through
a Commercially Proven Process. " Koppers Company,
Pittsburgh, Pennsylvania, August 1973.
(5) Leonard, Spike, Personal Communication. Koppers
Company, Pittsburgh, Pennsylvania, June 1974.
XV-26
-------
INTRODUCTION
The following comprehensive bibliography has been compiled to
assist persons researching the subject of clean fuels conversion pro-
cesses. This reference material is divided into five categories:
General
Process
Effluents
Economics
Miscellaneous
XVI-1
-------
REFERENCES - General
Blake, R. J. "How Acid-Gas Treating Processes Compare. " The Oil
and Gas Journal (January 9, 1967), 150-158.
Besselievre, E.B. The Treatment of Industrial Wastes. New York:
McGraw-Hill Book Co., 1969.
"Char Oil Energy Development. " U. S. Office .of Coal Research R&D
Report, No. 56, Final Report. Washington, D. C. : U.S. Government
Printing Office, 1971.
"Char Oil Energy Development. " U. S. Office of Coal Research R&D
Report, No. 73, Interim Report, No. 1. Washington, D. C.: U.S.
Government Printing Office, 1972.
Chopey, Nicholas P. "Coal Gasification: Can it Stage a Comeback?"
Chemical Engineering (April 3, 1972), 44+.
Davies, H.S., J.A. Lacey, andB.H. Thompson. "Processes for the
Manufacture of Natural Gas Substitutes. " J.I.G.E., Vol. 9 (1969), 375.
Dent, F. J., and D. Hebden. "The Gasification of Oil to Yield High
Calorific Value Gases. " World Power Conference Moscow 1968, .
Section A2, 207.
"Design and Operation for Air Pollution Control. " Paper presented at
the Metropolitan Engineers Council on Air Resources, October 24, 1968.
Dirksen, H.A., andH.R. Linden. Pipeline Gas from Coal by Methanation
of Synthesis Gas. Research Bulletin No. 31.- Chicago: Institute of Gas
Technology, July 1963.
"Electricity from Coal. " Mechanical Engineering (January 1959), 24+.
Elliott, M. A. "Research and Progess in the Production and Use of
Coal." Technical Paper No. 4. Washington, D. C. : National Resources
Planning Board, n. d.
Elliott, M.A., andH.R. Linden. "Gas, Manufactured." Kirk-Othmer
Encyclopedia of Chemical Technology, 2nd edition, Vol. 10, New York:
Interscience, 1966, 353-442.
"Energy Research Needs. " A Report by Resources for the Future, Inc.
for the National Science Foundation, October 1971.
XVI-2
-------
"Evaluation of Waste Waters From Petroleum and Coal Processing. "
Environmental Protection Agency report R2-72-001, December 1972.
Fitzgerald, K. H., and J.A. Richardson. "How Gas Composition
Affects Treating Process Selection. " Hydrocarbon Processing,
Vol. 45, No. 7 (July 1966), 125-129.
FMC Corporation, ""char Oil Energy Development." Research and
Development Report No. 56, Interim Report No. 1, prepared for the
Office of Coal Research, U.S. Department of the Interior. Washington,
D.C. : May 1970.
"Gasification: A Rediscovered Source of Clean Fuel. " Science, Vol. 178
(October 6, 1972), 44+.
Geer, M.R. Disposal of Solid Wastes from Coal Mining in Washington,
Oregon, Montana. U.S. Department of the Interior, Bureau of Mines,
Publication 8430, 1969.
Goar, E.G. "Today's Gas Treating Processes. " Oil and Gas Journal
(July 12, 1972 and July 19, 1972).
Goar, E.G. "Today's Sulfur Recovery Processes. " Hydrocarbon
Processing, Vol. 47, No. 9 (September 1968), 248-251.
Handbook of Chemistry and Physics. 44th edition. Cleveland, Ohio:
The Chemical Rubber Co.
i
Hoogendoorn, J. "17 Years of Gas Production From Coal. " Gas,
Vol. 33 (September 1972).
Johnson, W.B. "Coal Beats Oil Here. " Petroleum Refiner, Vol. 35
(December, 1956), 222-228.
Katel, S. "Removing Sulfur Dioxide From Flue Gases. " Chemical
Engineering Progress, Vol 62, No. 10 (October 1966), 67-73.
Kirk, et. al. Encyclopedia of Chemical Technology.
Kohl, A.L., andF.C. Riesenfeld. Gas Purification. New York:
McGraw-Hill, 1960, 346-369.
Kohl, A.L., andF.C. Riesenfeld. "Today's Processes for Gas
Purification." Chemical Engineering (June 15, 1959), 127-178.
XVI-3
-------
Lee, A. L., andH.L. Feldkirchner. "Methanation for Coal Hydrogasi-
fication. " Presented at the Joint Meeting of the Division of Fuel Chemistry
and the Division of Petroleum Chemistry of the American Chemical
Society, Chicago, September 9, 1970.
Linden, H.R. "Coal (Gasification). " Kirk-Othmer Encyclopedia of
Chemical Technology, 2nd Edition, Suppl. Vol. New York: Interscience,
1971, 198-217.
Linden, Henry R. "The Outlook for Synthetic Fuels. " Paper presented
to the Second Annual Meeting of API, Division of Paper Production,
Houston, March 6-8, 1972.
Linden, H.R. "A Reassessment of the Prospects for Coal Gasification. "
Coal Age, Vol. 76 (May 1971), 73-79.
Linden, H.R., and M.A. Elliott. "High-Btu Gases From Fluid Fuels-
Present and Future Methods. Part 2. Production of High-Btu Gas From
Natural Gas Liquids and Petroleum Oils. " American Gas Journal,
Vol. 186 (February 1959), 22-27.
Linden, H.R., J.J. Guyer, andE.S. Pettyjohn. "Production of Natural
Gas Substitutes by Pressure Hydrogasification of Oils. " A. G. A.
Processes (1954), 639-54.
Long, G. "Why Methanate SNG?" Hydrocarbon Processing, Vol. 51,
No. 8 (August 1972).
McWilliams, E.G., andR.P. Schuller. "SNG, Naphtha and Low -
Sulfur Fuel Oils from Crude Oils Using Commercially Proved Tech-
nology. " American Institute of Chemical Engineering (November to
December 1972).
Meyers, R.A. "Desulfurization of Coal. " Science. Vol. 177
(September 29, 1972), 1187+.
Mills, G.A. "Progress in Gasification--U. S. Bureau of Mines. "
Third American Gas Association Synthetic Pipeline Gas Symposium,
Chicago, November 17-18, 1970.
"This Month in Mining - Symposium Sees Hard Times Ahead for Industry
In Clean-Up of Tailing and Milling Wastes. " Engineering/Mining
Journal (May 1970), 102-104.
XVI-4
-------
Nelson, W.L. Petroleum Refinery Engineering. 4th Edition. New York:
McGraw-Hill Book Co., 1958.
Perry (ed.). Chemical Engineers' Handbook. New York: McGraw-Hill.
Ray. "Oil Refining Wastes. " Sewage Industrial Wastes (November 1958).
Ricketts, T.S., and D. C. Elgin. "Proceedings at the Joint Conference
on Gasification Processes.11 London: Institute of Fuel, 1962, C-l.
Ruggles. "Basic Petrochemical Processes as Waste Sources. "
Sewage Industrial Wastes (March 1959).
Seglin, L., and. R. T. Eddinger. "Coal (to Synthetic Crude Oil). "
Kirk-Othmer Encyclopedia of Chemical Technology. 2nd Edition,
Suppl. Vol. New York: Interscience, 1971, 178-198.
Smith, and Stinton. Fuels and Combustion. New York: McGraw-Hill,
1952.
"SNG From Hydrocarbon Liquids. " From SNG Symposium I sponsored
by IGT, March 12, 1973.
Squires, Arthur M. "Clean Power From Coal. " Science, Vol. 169,
No. 3948 (August 28, 1970).
"The Supply-Technical Advisory Task Force — Synthetic Gas-Coal. "
Report for the Federal Power Commission by the Synthetic Gas-Coal
Task Force, April 1973.
"Synthetic Fuels: What, Which?" Chemical Engineering (April 17,
1972), 62+.
Thompson, B.H., andH.L. Conway. "The Hydrogenation of Hydro-
carbons in Relation to the Manufacture of Substitute Natural Gas. "
Paper presented to the 65th Annual Meeting of the American Institute
of Chemical Engineering.
Thompson, B.H., B.B. Majumdar, andH.L. Conway. "The Hydro-
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TECHNICAL REPORT DATA
(Please read Instructions on the reverse before completing)
1. REPORT NO.
EPA-450/3-75-028
2.
4. TITLE AND SUBTITLE
Emissions from Processes Producing Clean Fuels
3. RECIPIENT'S ACCESSION-NO.
5. REPORT DATE
March 1974-Date of issue
6. PERFORMING ORGANIZATION CODE
7. AUTHOR(S)
F. Glazer, A. Hershaft, & R. Shaw
8. PERFORMING ORGANIZATION REPORT NO.
9075--015
9. PERFORMING ORG-VNIZATION NAME AND ADDRESS
Booz-Allen & Hamilton, Inc.
4733 Bethesda Avenue
Bethesda, Md. 20014
10. PROGRAM ELEMENT NO.
11. CONTRACT/GRANT NO.
68-02-1358
12. SPONSORING AGENCY NAME AND ADDRESS
Environmental Protection Agency
Office of Air Quality Planning & Standards
Research Triangle Park,
North Carolina 27711
13. TYPE OF REPORT AND PERIOD COVERED
Final
14. SPONSORING AGENCY CODE
15. SUPPLEMENTARY NOTES
16. ABSTRACT
Processes for converting coal, oil and shale to clean fuels were studied. Most
of the processes considered have not yet been developed to commercialization. Based
on preliminary data and engineering estimates developed, it can be concluded that:
(1) Unit processes apparently exist to provide adequate control of most
waste streams from clean-fuel plants, although some pollutants remain
difficult to treat;
(2) The pollutant emissions from clean-fuel processes will, in general,
be relatively low when compared to emissions from alternative uses
of the raw fuel;
(3) The total costs of pollutant control (1973 basis) for achieving low
levels of plant emissions, in high-Btu coal gasification processes,
range from 15 to 30 cents per million Btu and can be expected to
increase in proportion to the severity of the emission controls required;
(4) Low-Btu gasification followed by desulfurization may be considered as
a pollution control system because it permits the combustion of coal
with decreased emissions. In that case the overall cost of pollution
control is about $Q.65/1Q6 Btu of gas produced. (Bibliography included.)
17.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b.lDENTIFIERS/OPEN ENDED TERMS
c. COSATI Field/Group
Fuels
Coal gasification
Air Pollution
Desulfurization
Fuel conversion
Sulfurous emissions
21/D
13/H
13/B
7/A
18. DISTRIBUTION STATEMENT
Unlimited
19. SECURITY CLASS (This Report)'
Unclassified
21. NO. OF PAGES
4*4
20. SECURITY CLASS (Thispage)
Unclassified
22. PRICE
EPA Form 2220-1 (9-73)
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