EPA 600/2-76-008
January 1976
Environmental Protection Technology Series
SO, CONTROL PROCESSES FOR
2
NON-FERROUS SMELTERS
Industrial Environmental Research Laboratory
Office of Research and Development
U.S. Environmental Protection Agency
Research Triangle Park, North Carolina 27711
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RESEARCH REPORTING SERIES
Research reports of the Office of Research-and Development,
U.S. Environmental Protection'Agency, have been grouped into
five series. These five broad categories were established to
facilitate further development and application of environmental
technology. Elimination of traditional grouping was consciously
planned to foster technology transfer and a maximum interface in
related fields. The five series are:
1. Environmental Health Effects Research
2. Environmental Protection Technology
3. Ecological Research
4. Environmental Monitoring
5. Socioeconomic Environmental Studies
This report has been assigned to the ENVIRONMENTAL PROTECTION
TECHNOLOGY series. This series describes research performed
to develop and demonstrate instrumentation, equipment and
methodology to repair or prevent environmental degradation from
point and non-point sources of pollution. This work provides the
new or improved technology required for the control and treatment
of pollution sources to meet environmental quality standards.
EPA REVIEW NOTICE
This report has been reviewed by the U.S. Environmental Protection
Agency, and approved for publication. Approval does not signify that
the contents necessarily reflect the views and policies of the Agency, nor
does mention of trade names or commercial products constitute endorse-
ment or recommendation for use.
This document is available to the public through the National
Technical Information Service, Springfield, Virginia 22161.
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EPA-600/2-76-008
SO2 CONTROL PROCESSES
FOR
NON-FERROUS SMELTERS
by
John C. Mathews, Faust L. Bellegia,
Charles H. Gooding, and George E. Weant
Research Triangle Institute
P.O. Box 12194
Research Triangle Park, NC 27709
Contract No. 68-02-1491
ROAP No. 21ADC-059 and 21AUY-040 and -041
Program Element No. 1AB013 and 1AB015
EPA Project Officer: Douglas A. Kemnitz
Industrial Environmental Research Laboratory
Office of Energy, Minerals, and Industry
Research Triangle Park, NC 27711
Prepared for
U.S. ENVIRONMENTAL PROTECTION AGENCY
Office of Research and Development
Washington, DC 20460
January 1976
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ABSTRACT
This report provides a review and evaluation of a number of
absorption-based SCL control systems and the application of these
control systems to those U.S. primary copper smelters which generate a
weak SCL-containing gas stream.
Capital and operating cost relationships have been developed for
each specific process covering a range of gas flows and SO- concentrations.
Separate general costs for gas pretreatment and the end-of-the-line S0»
utilization facilities, i.e., sulfuric acid, elemental sulfur, and
liquid SO- plants, have also been provided.
Those thirteen U.S. primary copper smelters which still generate
weak SCv streams have been reviewed with reference to their current
operation and active programs in hand to control or eliminate weak SO^
streams. Appropriate SO™ control processes have been matched with the
individual smelters, and related capital and operating costs were developed
from the earlier established cost relationships.
ii
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S02 CONTROL PROCESSES FOR NON-FERROUS SMELTERS
TABLE OF CONTENTS
Abstract ii
List of Figures . v
List of Tables i viii
Units of Measure - Conversions xii
Acknowledgments xiii
Sections
1.0 Summary . . .' 1
2.0 General Discussion 3
3.0 Sulfuric Acid 19
4.0 Lime/Limestone Scrubbing 41
5.0 Sodium Scrubbing-Regenerative (Wellman-Lord) 59
6.0 Double-Alkali Sodium Base (Throwaway Process) .... 75
7.0 Magnesium Oxide Scrubbing 95
8.0 Dimethylaniline/Xylidine Process 115
9.0 Citrate Process 129
10.0 Ammonia Process .,.*.. 149
11.0 Application of Absorption-Based S02 Control Systems to
Weak S02 Copper Smelter Reverberatory Furnaces ... 177
12.0 Application of Absorption-Based S02 Control Systems to
the Weak S02 Gas Streams of U.S. Primary Copper
Smelters ....... 183
12.1 The Anaconda Company - Anaconda, Montana . . . 183
12.2 ASARCO - El Paso Smelter - El Paso, Texas . . . 185
12.3 ASARCO - Hayden Smelter - Hayden, Arizona . . . 197
12.4 ASARCO - Tacoma, Washington Smelter - Tacoma
Washington 205
12.5 Kennecott Copper Corporation - Garfield, Utah - 215
12.6 Kennecott Copper Corporation - Hayden, Arizona. 217
12.7 Kennecott Copper Corporation - Hurley, New Mexico 225
12.8 Kennecott Copper Corporation - McGill, Nevada . 233
12.9 Magma Copper Company - San Manuel, Arizona . . 241
12.10 Phelps Dodge Corporation - Ajo, Arizona ....
12.11 Phelps Dodge Corporation - Douglas, Arizona . . 257
12.12 Phelps Dodge Corporation - Morenci, Arizona . . 265
12.13 White Pine Copper Company - White Pine, Michigan 273
iii
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Table of Contents (cont'd)
Page
Appendices
Gas Cleaning and Conditioning 281
S02 Absorption Costs 289
Auxiliary Sulfuric -Acid Plant 292
Auxiliary Sulfur Plant 298
SCv Liquefaction and Storage Costs 302
Utility Costs 304
Capital and Annual Operating Cost Computation Sheets for
Selected S02 Control Processes Matched to Specific Primary
Copper Smelters 308
iv
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LIST OF FIGURES
Figure
3-1
3-2
3-3
3-4
3-5
3-6
3-7
3-8
3-9
3-10
3-11
4-1
4-2
4-3
4-4
5-1
5-2
5-3
5-4
6-1
6-2
i
6-3
6-4
7-1
7-2
7-3
7-4
Total Capital Investment Costs
Total Capital Investment Costs, Acid Plant Section
Total Capital Investment Costs, Gas Cleaning Section
Total Capital Investment Costs, Refrigeration Section
Total Capital Investment Costs, Preheating Section
Total Annual Direct Operating Costs, Acid Section . .
Total Annual Direct Operating Costs, Gas Cleaning
Total Annual Direct Operating Costs, Refrigeration
Section
Total Annual Direct Operating Costs, Preheating
Total Annual Direct Operating Costs
Total Capital Investment Costs, Limestone Handling
and Disposal •
Wellman-Lord Process, Regenerative Sodium Scrubbing .
Total Capital Investment Costs . .
Total Capital Investment Costs, S02 Regeneration
Section
Sodium Base Double Alkali Process-Dilute System . . .
Total Capital Investment Costs
Total Capital Investment Costs Regeneration & Pre-
cipitate Handling .
Total Capital Investment Costs, S02 Regeneration
Section
Page
21
30
31
32
33
34
35
36
37
38
39
43
54
55
56
61
69
70
71
77
89
90
91
97
110
111
112
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List of Figures (cont'd)
Figure Page
8-1 Dimethylaniline/Xylidine Process 117
8-2 Total Capital Investment Costs 125
8-3 Total Annual Direct Operating Costs . 126
9-1 Citrate Process 131
9-2 Total Capital Investment Costs 143
9-3 Total Annual Direct Operating Costs 144
9-4 Total Capital Investment Costs, SOg Regeneration
Section 145
10-1 Ammonia Scrubbing Process-ABS Acidulation 151
10-2 Total Capital Investment Costs 171
10-3 Total Annual Direct Operating Costs 172
10-4 Total Capital Investment Costs, S02 Regeneration
Section 173
10-5 Total Capital Investment Costs, 4-S.tage Ammoniacal
Scrubber with Liquor Interceding 174
12.2-1 ASARCO-E1 Paso, Texas-Smelter Flow Schematic .... 187
12.3-1 ASARCO-Hayden, Arizona-Smelter Flow Schematic . . . 199
12.4-1 ASARCO-Tacoma, Washington-Smelter Flow Schematic . . 207
12.6-1 Kennecott Copper-Hayden, Arizona-Smelter Flow
Schematic 219
12.7-1 Kennecott Copper-Hurley, New Mexico-Smelter Flow
Schematic 227
12.8-1 Kennecott Copper-McGill, Nevada-Smelter Flow
Schematic 235
12.9-1 Magma Copper Company-San Manuel, Arizona-Smelter
Flow Schematic 243
12.10-1 Phelps Dodge-Ajo, Arizona-Smelter Flow Schematic . . 253
12.11-1 Phelps Dodge-Douglas, Arizona-Smelter Flow Schematic 259
12.12-1 Phelps Dodge-Morenci, Arizona-Smelter Flow Schematic 267
12.13-1 White Pine Copper Company-White Pine, Michigan-
Smelter Flow Schematic 275
A-l Gas Cooling & Conditioning for Regenerative S02
Absorption Systems 283
A-2 Total Capital Investment Costs & Total Annual
Direct Operating Costs 286
vi
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List of Figures (cont'd)
Figure
B-l S02 Absorption Systems, Total Capital Investment
Costs 291
C-l Auxiliary Sulfuric Acid Plant, Total Capital
Investment Costs & Total Annual Direct Operating
Costs' (Based on 8-9% S02 to Acid Plant) 295
D-l Auxiliary Sulfur Plant, Total Capital Investment
Costs & Total Annual Direct Operating Costs ..... 300
E-l Liquefaction and Storage of S02 303
F-l Capital Cost of Water Treating Facilities 305
F-2 Capital Cost of Package Boilers 306
F-3 Capital Cost of Electrical Substations 307
vii
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LIST OF TABLES
Table Page
2.1 Status of S02 Control Processes 6
2.2 Effect of Byproduct Credit on Net Total Annualized
Costs of the Regenerable SC^ Control Processes
When Treating a Gas Stream Containing 1% S02 ... 16
3.1 Sulfuric Acid, Total System Capital and Total
Annual Costs 40
4.1 Lime/Limestone Process, Unit Usage and Cost Data . 53
4.2 Lime/Limestone Process, Capital and Total Annual
Costs 57
4.3 Lime/Limestone Process, Total System Capital and
Annual Operating Costs 58
5.1 Sodium Scrubbing Regenerative Wellman-Lord Process,
Unit Usage and Cost Data 68
5.2 Sodium Scrubbing Regenerative Wellman-Lord Process,
Capital and Total Annual Operating Costs 72
•
5.3 Sodium Scrubbing Regenerative Wellman-Lord Process,
Total System Capital and Annual Operating Costs . . 73
6.1 Sodium Scrubbing-Double Alkali (Throwaway), Unit
Usage and Cost Data 88
6.2 Sodium Scrubbing-Double Alkali Dilute System (Throw-
away), Capital and Total Annual Cost 92
6.3 Sodium Scrubbing-Double Alkali Dilute System (Throw-
away), Total System Capital and Annual Operating Costs 93
7.1 Magnesium Oxide Process Unit Usage and Cost Data . 109
7.2 Magnesium Oxide Scrubbing, Capital and Total Annual
Costs 113
7.3 Magnesium Oxide Scrubbing, Total System Capital and
Annual Costs 114
8.1 DMA/Xylidine Process, Unit Usage and Cost Data . . 124
8.2 DMA/Xylidine Process, Capital and Total Annual
Operating Costs 127
8.3 DMA/Xylidine Process, Total System Capital and
Annual Operating Costs 128
9.1 Citrate Process Unit Usage and Cost Data 142
9.2 Citrate Process, Capital and Total Annual Costs . . 146
9.3 Citrate Process, Total System Capital and Annual
Operating Costs 147
10.1 Ammonia Process Unit Usage and Cost Data 170
viii
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List of Tables (cont'd)
Table
10.2 Ammonia Process, Capital and Total Annual Costs . . . 175
10.3 Ammonia Process, Total System Capital and Annual
Costs 176
12.2-1 ASARCO-E1 Paso Smelter, Texas-Characterization of
Weak SC>2 Gas Streams Based on Present Operation . . . 188
12.2-2 ASARCO-E1 Paso Smelter, Texas-Characterization of
Weak S02 Gas Streams Adjusted for Concentrate Change 192
12.2-3 ASARCO-E1 Paso-Capital Costs for S02 Control Processes
on Combined Roaster and Reverberatory Furnace
Off-Gases 193
12.2-4 ASARCO-El Paso-Annual Operating Costs and Total
Net Annualized Costs for S02 Control Processes on
Combined Roaster and Reverberatory Furnace Off-Gases 194
12.2-5 ASARCO-El Paso-Capital Costs for SO, Control
Processes on Reverberatory Furnace Off-Gases Only . . 195
12.2-6 ASARCO-El Paso-Annual Operating Costs and Total Net
Annualized Costs for SO2 Control Processes On Rever-
beratory Furnace Off-Gases Only 196
12.3-1 ASARCO-Hayden, Arizona-Characterization of Weak
S02 Gas Streams 202
12.3-2 ASARCO-Hayden-Capital Costs for S02 Control Processes
on Combined Roaster and Reverberatory Furnace Off-
Gases 203
12.3-3 ASARCO-Hayden-Annual Operating Costs and Total Net
Annualized Costs 204
12.4-1 ASARCO-Tacoma, Washington-Characterization of Weak
S02 Gas Streams 211
12.4-2 ASARCO-tacoma-Capital Costs for S02 Control Processes
on Combined Roaster and Reverberatory Furnace Off-
Gases • 212
12.4-3 ASARCO-Tacoma-Annual Operating Costs and Total Net
Annualized Costs for S02 Control Processes on Com-
bined Roaster and Reverberatory Furnace Off-Gases . . 213
12.6-1 Kennecott Copper Corp.-Hayden, Arizona-Characteri-
zation of Weak S0~ Gas Streams 222
12.6-2 Kennecott Copper Corp.-Hayden-Capital Costs for S02
Control Processes on Reverberatory Furnace Off-Gases 223
12.6-3 Kennecott Copper Corp.-Hayden-Annual Operating Costs
and Total Net Annualized Costs 224
ix
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List of Tables (cont'd)
Table Page
12.7-1 Kennecott Copper-Hurley, New Mexico-Characterization
of Weak S02 Gas Streams 230
12.7-2 Kennecott Copper-Hurley-Capital Costs for SC>2 Control
Processes on Reverberatory Furnace Off-Gases .... 231
12.7-3 Kennecott Copper-Hurley-Annual Operating Costs and
Total Net Annualized Costs 232
12.8-1 Kennecott Copper- McGill, Nevada-Characterization
of Weak S02 Gas Streams 238
12.8-2 Kennecott Copper-McGill-Capital Costs for S02 Control
Processes on Reverberatory Furnace Off-Gases .... 239
12.8-3 Kennecott Copper-McGi11-Annual Operating Costs and
Total Net Annualized Costs 240
12.9-1 Magma Copper-San Manuel, Arizona-Characterization of
Weak S02 Gas Streams 247
12.9-2 Magma Copper-San Manuel-Capital Costs for S02 Control
Processes on Reverberatory Furnace Off-Gases .... 248
12.9-3 Magma Copper-San Manuel-Annual Operating Costs and
Total Net Annualized Costs 249
12.10-1 Phelps Dodge-Ajo, Arizona-Characterization of Weak
S02 Gas Streams 254
12.10-2 Phelps Dodge-Ajo-Capital and Annual Operating Costs
for DMA/Xylidine Control System on Reverberatory
Furnace Off-Gases 255
12.11-1 Phelps Dodge-Douglas, Arizona-Characterization of
Weak S02 Gas Streams 262
12.11-2 Phelps Dodge-Douglas-Capital Costs for S02 Control
Processes on Weak S02 Smelter Off-Gases 263
12.11-3 Phelps Dodge-Douglas-Annual Operating Costs and Total
Net Annualized Costs 264
12.12-1 Phelps Dodge-Morenci, Arizona-Characterization of
Weak S02 Gas Streams 269
12.12-2 Phelps Dodge-Morenci-Capital Costs for S02 Control
Processes on Reverberatory Furnace Off-Gases .... 270
12.12-3 Phelps Dodge-Morenci-Annual Operating Costs and Total
Net Annualized Costs 271
12.13-1 White Pine Copper-White Pine, Michigan-Characteriza-
tion of S02 Gas Streams 278
12.13-2 White Pine Copper-White Pine-Capital Costs for S02
Control Processes on Converter Off-Gases 279
12.13-3 White Pine Copper-White Pine-Annual Operating Costs
and Total Net Annualized Costs . . 280
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List of Tables (cont'd)
Gas Cleaning and Conditioning Unit Usage and
Cost Data 285
A-2 Gas Cleaning and Conditioning for Regenerable S02
Control Processes Capital and Total Annual Costs . . 287
A-3 Gas Cleaning and Conditioning for "Throwaway" SC^
Control Processes Capital and Total Annual Costs . . 288
C-l Auxiliary Sulfuric Acid Plant Unit Usage and Cost
Data (8-9% S02 in Feed Gas) 294
C-2 Auxiliary Sulfuric Acid Plant (No gas conditioning)
Capital and Total Annual Costs 296
C-3 Auxiliary Sulfuric Acid Plant (Including Dry Gas
Cleaning) Capital and Total Annual Costs 297
D-l Auxiliary Sulfur Plant Unit Usage and Cost Data . . . 299
D-2 Auxiliary Elemental Sulfur Plant Capital and Total
Annual Costs • . . . . 301
xi
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UNITS OF MEASURE - CONVERSIONS
Environmental Protection Agency policy is to express all measure-
ments in Agency documents in metric units. When implementing this
practice will result in undue cost or lack of clarity, conversion
factors are provided for the non-metric units used in a report.
Generally, this report uses British units of measure. For conversion
to the metric system, use the following conversions:
To convert from
cfm
°F
ft.
gal.
gpm
gr/scf
hp
in.
in. we
Ib
Ib/hr
To
m /sec
°C
m
1
I/sec
mg/Nm
W
m
N/m2
kg
kg/hr
Multiply by
0.0004719
5/9 (°F-32)
0.3048
3.785
0.0631
2288.136
745.7
0.0254
248.84
0.454
0.454
xii
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ACKNOWLEDGMENTS
The authors wish to acknowledge the assistance provided during the
course of the study by Control Systems Laboratory personnel, staff
members of EPA Regions 6, 8 and 9, and representatives of the
responsible State air pollution agencies.
We also wish to recognize the cooperation and help extended by
the representatives of the different copper smelters contacted during
the second phase of the program.
xiii
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1.0 SUMMARY
This study has been directed towards the review and evaluation of
certain specified scrubbing or absorption-based SO- control systems and
the application of these control systems to those U.S. primary copper
smelters which generate weak SCL-containing gas streams. The number of
absorption-based systems theoretically available today for flue gas
desulfurization is large. Many of these systems have had limited
application at the pilot plant level under specialized conditions and
their general capability is uncertain. The scope of this study has been
limited to those processes which have been or are being extensively
evaluated in the U.S. and whose feasibility has been demonstrated at
pilot plant level under semi-commercial plant conditions or with full
scale installations. It is recognized that other processes such as the
Chioyda Thoroughbred and the Powerclaus (phosphate scrubbing) processes
are also available and appear to have considerable promise, but it is
believed that the processes selected for study represent the range and
feasibility of present flue gas desulfurization approaches.
Each selected control process has been given a broad overview with
particular reference to operational considerations, development status,
desulfurization efficiency, oxidation, gas pretreatment requirements,
energy demand and cost.
Capital and operating cost relationships in terms of gas flow and
SC>2 content have been developed for grass-roots installation of each
specific process together with general costs for gas pretreatment and
end-of-the-line S02 utilization facilities such as sulfuric acid, ele-
mental sulfur and liquid SO™ plants.
Of the fifteen U.S. primary copper smelters, two facilities—Cities
Service, Copperhill, Tennessee and Inspiration Copper at Miami, Arizona-
utilize metallurgical processes which do not produce dilute S02 content
streams; and these two smelters, accordingly, have not been included in
this study. Several of the remaining smelters have active programs in
hand to eliminate the generation of weak SO^-containing gas streams by
the use of new or developing technology. These particular smelters and
their programs have been noted.
A study attempting to apply S02 control systems which generally
have had only preliminary commercial application in the specialized
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industrial area of power utilities to the special situation of primary
copper smelters faces a number of difficulties. There are the technical
and economic uncertainties associated with any projection, particularly
a projection which must make use of limited application data and cost
information, frequently only broadly defined. Perhaps an even more
difficult area to respond to involves the final disposition or utilization
of the recovered sulfur dioxide. In the real world this consideration
cannot be divorced from the determination of the total costs of any SO-
control system. Sulfur dioxide recovered from a smelting process may be
conveniently and readily used in an existing or new sulfuric acid plant,
but unless the resulting acid can be continuously marketed even at loss
or break-even conditions, neutralization facilities with the concomitant
problem of solids disposal become an essential complementary part of the
SO™ control system with a significant contribution in both capital and
operating costs.
This study, under its defined scope, does not include either marketing
or special disposal considerations. Control processes have been matched
with the individual smelters on the basis of the expected response of
the control process itself to the conditions existing at that smelter
and the characteristics of that smelter's weak S02 streams. It is also
emphasized that no consideration has been given to the possibility of
minimizing air dilution in the offgas flue system and thereby reducing
the costs of gas cleaning and S02~absorption sections of any control
system. Equipment and flue conditions and operating modes of each
individual smelter will be principal factors in determining how much air
dilution can be reduced and at what cost. The scope of the program
does not include this evaluation. Therefore the resulting cost profile,
both capital and operating, is not a total confirmation of the cost of
providing an overall S02 control system, but rather an order-of-magnitude
indication of the "grass-roots" costs of installing a particular SO-
control system. The final "turn-key" cost of this system can be only
established by an in-depth analysis of the specific smelter's operation
as it affects the weak gas system and by evaluating the potential of in-
house usage if sulfuric acid is produced, local area marketing prospects,
or the impact of providing acid neutralization facilities.
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2.0 GENERAL DISCUSSION
2.1 BACKGROUND
Although the non-ferrous smelting industry is the second largest
group emitter of S02 in the United States, attention over the past 5 to
7 years has focused on the utility industry with its 65-percent share of
the total U.S. sulfur emissions. Process developers and engineers,
responding to the demand for viable processes to reduce the level of S02
in utility flue gases, have directed their attention to the special
problems posed by the very large gas volumes and very low S02 concen-
trations associated with these gas streams. Thus, in the United States
over the past 4 to 5 years, process development, pilot plant installations,
and commercial demonstration units have been almost exclusively devoted
to S02 control processes applied to utility flue gases. The result
today is that there exists a considerable amount of process, engineering,
and cost data on S02 control methods applied to utility flue gases at
both the pilot plant and commercial demonstration levels, and there is a
corresponding degree of confidence in projecting expected performance
levels and capital investment requirements.
In the non-ferrous smelter area, some limited pilot plant work on
weak S02 gas streams with the same S02 control processes as being applied
to utilities has been conducted or is planned for the immediate future
at a number of smelter locations. However, cost estimates for full
scale installations are limited and tend to indicate "range of costs"
only.
* - . . . . ...•.,• . , . -
a) 4000 SCFM plant on copper reverberatory furnace gases at
McGill, Nevada Smelter of Kennecott Copper Corp. to evaluate primary
limestone scrubbing (1972); b) 300 SCFM plant on copper reverberatory
furnace gas at San Manuel, Arizona Smelter of Magma Copper Co. to evaluate
Bureau of Mines citrate process (1970-71); c) 1000 SCFM plant on Kellogg,
Idaho lead smelter of Bunker Hill Co. to investigate citrate process.
Part of this system started up in Feb. 1974-continuing; and d) 4000*SCFM
plant on copper reverberatory furnace gas started up April 1975 at San
Manuel, Arizona Smelter of Magma Copper Co. to investigate ammonia
double alkali system and others.
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In evaluating the capabilities of the specified SO- control pro-
cesses, the literature and published reports of recent or on-going
development work in the utility area provide sufficient input to establish
a qualitative assessment of performance under smelter conditions.
However, the development of both capital and operating costs for the gas
flows and SO- concentrations which might be expected in non-ferrous
smelters requires considerable massaging, and in certain instances
adjustment of the available utility-oriented cost data to ensure that
the determined cost estimates for the different processes as applied to
smelter operations are reasonably compatible.
2.2 CHARACTERISTICS OF WEAK S02 SMELTER GASES
If weak S02~containing gases are defined as those containing less
than 4 percent SO-, the source of such streams in primary non-ferrous
smelters is usually:
1) the multi-hearth roasters in primary copper smelters
2) the reverberatory furnaces in primary copper smelters
3) the sintering machine and blast furnace in lead smelters
4) the Ropp roaster (if used) and sintering machine in zinc
smelters.
The flue gases in all cases leave the furnace at temperatures of
1200 to 1800°F (650°-982°C) with S02 concentrations which may range from
0.5 percent to 2.5 percent and with considerable amounts of mechanically
entrained solids and vaporized metal compounds. The quantity and nature
of this material depends on the operating conditions of the furnace—
charging method, gas velocity, temperature—and the composition of the
charge itself. Volatile compounds of such elements as arsenic, antimony,
bismuth, cadmium, mercury, selenium, tellurium and thallium, as well as
of copper, lead and zinc, may be present. SO- is also formed during the
smelting process with the amount formed related to the amount of excess
oxygen and the presence of metal oxides such as iron oxide which may
catalyze the oxidation of S02 to SO-.
Under usual smelter practice, it is common for the heat value of
the flue gases from reverberatory furnaces to be recovered via a waste
heat boiler and most of the entrained dust and fume recovered for its
economic value in dry-type particulate removal devices such as balloon
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flues, cyclones and electrostatic precipitators before the gas stream is
vented to the atmosphere at temperatures of around 400°F (204°C). Thus
the exit gas feed available to any S02 recovery process is at an elevated
temperature; it contains residual quantities of solid and fume particulates;
and in addition to S02 and S0», it may contain gaseous contaminants such
as HC1 and HF. The limitations of most S02 absorption-based control
processes to high temperatures and to oxidation side-reactions make
pretreatment or conditioning of this gas stream an essential step in the
application of such processes.
Under direction of the Project Officer, this study was addressed
directly to the weak gas streams from copper smelters, i.e., reverberatory
furnaces and multi-hearth roasters. To provide a standardized approach
for each SO^ control process under review, the gas feed available at the
battery limits of the SCL control section has been specified as being
cooled to 600°F (316°C) by the use of radiation cooling and waste heat
boilers, and the particulate loading has been reduced to approximately
0-1 gr/SCF via cyclones and electrostatic precipitators.
2.3 S02 CONTROL PROCESSES
The S02 control processes identified for evaluation and applicability
to the weak S02~containing effluent streams generated in non-ferrous
smelters represent a wide range in the level of demonstrated application,
reliability and control effectiveness.
These processes are:
1) sulfuric acid
2) lime and limestone scrubbing
3) sodium scrubbing regenerable (Wellman-Lord Process)
4) sodium scrubbing non-regenerable (Double-Alkali Process)
5) magnesium oxide
6) dimethylaniline (DMA.)/xylidine
7) sodium citrate
8) ammonia.
Table 2-1 provides a broad summary of their current states of development
and their main areas of application. With the exception of the sulfuric
acid process, these S02 control systems all have the same functional
processing steps:
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Table 2.1.
STATUS OF SO2 CONTROL PROCESSES
Process
Developmental Stage
Area of Application
Sulfuric Acid
Well understood and proven technology,
Widely used in non-ferrous smelters
or gas streams containing above 3,5%
S02- Can operate on weaker streams
with appropriate modifications.
Lime and Limestone Scrubbing
Commercial sized systems have been
installed and are being installed.
Considerable development work done
over past 5 to 6 years. Reliability
of operation improving rapidly.
Primarily in the power plant area,
but some metallurgical applications.
No known installations in non-ferrous
smelters.
Sodium Scrubbing Regenerable
(Wellman-Lord Process)
A number of commercial installations
in United States and Japan since
1970. Good demonstrated reliability.
Sulfuric acid plant and Glaus plant
tail gas scrubbing and flue gas
from steam generators. No known
smelter applications.
Sodium Scrubbing Non-
Regenerable (Double
Alkali Process)
Extensive development work in both
United States and Japan. Some
commercial applications in both
countries.
Power plant and industrial boiler
applications. No smelter installations.
Magnesium Oxide
Several commercial sized installations
in both United States and Japan.
Long-term reliability still to be
demonstrated.
In the United States, on utility power
plants; but in Japan, installed on tail
gas from sulfuric acid and Claus plant.
Two units installed in copper smelters.
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Table 2.1 (cont'd)
Process
Developmental Stage
Area of Application
Dimethylaniline (and Xylidine)
Process has been available for many
years and several commercial
installations are in operation in
United States.
Several DMA installations in
non-ferrous smelters handling a
range of S02 concentrations. Favors
higher concentratins. No xylidine
application in the United States.
Sodium Citrate
Second generation process still in
small scale pilot plant stage.
Pilot plant installations have
treated boiler flue gases, copper
reverberatory flue gases and zinc
sinter furnace gases, all with
high S02 removal efficiencies.
Ammonia (Regenerable)
Process has been investigated
extensively and commercial
applications in both Canada and the
USSR date back many years. Pilot
plant investigations current in
the United States are concentrating
on improved regeneration route.
There has been long term application
treating smelter offgases
(Consolidated Mining and Smelting Co.
of Canada) but effective regeneration
with reduced energy requirements
still to be commercially demonstrated.
-------
1) Gas cooling and conditioning
2) Absorption of the SCL from the gas stream
3) Regeneration (or disposal) of the SCL from the absorbent
liquid
4) Utilization of the regenerated SCL in regenerable systems
via sulfuric acid or elemental sulfur plants.
The steps of gas cooling and conditioning, and S02 utilization can be
defined independently of the absorption and regeneration (or disposal)
steps which may be considered to constitute the basic SCL control process
itself.
2.4 GAS COOLING AND CONDITIONING
The purpose of this step is:
a) to cool and humidify the gas stream to prevent excessive
evaporation and deposition of solids in the SO™ absorber
itself
b) to remove the residual particulate matter which may catalyze
undesirable side reactions in the absorber or contribute to
high rates of build-up of inert material in closed loop or
regenerative systems.
The degree of pretreatment required by the gas feed to an S02 absorption-
based system will thus depend on both the characteristics of the gas
stream itself and the limitations of the SO- control process. The
control process may also provide options which allow some economic
trade-offs between higher conditioning capital investment and modified
operational modes downstream, although with mechanically precleaned
smelter gases and the small particulate size of the residual material,
the potential may be limited.
Specific industrial experience on S02 control systems is limited,
and largely related to the conditioning requirements of utility flue
gases on throwaway lime/limestone based systems. However, a number of
sulfuric acid plants have long operated with gas cooling and conditioning
systems under conditions similar to those existing in non-ferrous smelter
operations. J. M. Connor in a paper in 1968 provided discussion and
cost estimates for a typical sulfuric acid plant gas cooling and conditioning
system, and his costs, with appropriate modification and escalation,
2 3
appear to be reflected in several later studies. '
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A typical sulfuric acid plant conditioning system is usually made
up of three separate steps:
1) quenching the gas stream with water in a suitable contacting
device to provide essentially adiabatic cooling to saturation
together with some degree of particulate removal. Additional
cooling may be provided in either a second tower or by
recirculating the quench water through a cooling loop
2) removal of acid mist carryover with additional particulate
removal
3) neutralization and solids removal to waste of part of the
recirculated quenching water.
Step 2 is an important part of the overall conditioning process, and the
final specifications of the cleaned gas are determined largely by the
approach taken in this section. Multi-stage wet electrostatic pre-
cipitators in both series and parallel configurations are commonly used
in acid plants installed on smelter SO,, gas streams. In the Miami
Smelter of the Inspiration Consolidated Copper Company, converter and
electric furnace offgases feeding a new double absorption sulfuric acid
plant are cleaned and conditioned via venturi scrubbers and packed
washing towers, followed by two stages of electrostatic mist precipitators
with four separators in parallel in each stage. Capital costs for such
cleaning systems may constitute as much as 40 to 45 percent of the total
investment for a sulfuric acid plant installation.
The degree of gas cleaning required for SO™ control processes
varies widely depending on the characteristics of the process itself.
Conditioning requirements for "throwaway" processes such as the lime/
i
limestone process are essentially satisfied by providing gas cooling and
humidification only. Regenerable processes, however, may require additional
feed gas cooling and/or a higher degree of partieulate and acid mist
removal, although it is unlikely that any of these processes would
require full "acid plant level" pretreatment.
Capital and operating costs have been developed in Appendix A for
two separate gas cooling and conditioning systems which would be expected
to satisfy the requirements of "throwaway" SO,, control processes and all
regenerable SO^ control processes, respectivelyt As noted above some
regenerable S02 control processes may allow process approaches in the
regeneration section which provide for the removal of gas-introduced
particulate material and thereby reduce the degree of initial feed gas
-------
cleaning needed. The viability of this approach is related to the size
distribution of the particulate material entering the system and the
operating characteristics and nature of the SO^ absorption column
itself. Non-ferrous gases which have passed through balloon flues,
waste heat boilers and electrostatic precipitators prior to the S02
control system will have a size distribution concentrated in the submicron
range, and it is unlikely that the low pressure drop absorption column
will exercise much influence on further particulate removal. However,
in any closed system, inert material will build up from make-up water,
from reagent impurities, and from the inevitable oxidation reactions
which can at best be controlled but not entirely eliminated; and these
materials must be removed or purged from the system. This purge will at
the same time tend to control any buildup of gas-introduced particulates.
If the active reagent is expensive, e.g., sodium salts in the Wellman-
Lord process or magnesium in the magnesium oxide process, direct purging
to waste may be unattractive and may suggest a secondary purge treatment
system to recover the active reagents, and sized to control particulate
build-up at the same time.
2.5 S02 ABSORPTION AND REGENERATION
In the SO,, absorption section, the cooled and cleaned S09-containing
£> b>
gas is scrubbed, usually countercurrently, with a particular absorbent
solution. The process can be carried out using a packed tower, sieve
plate or impingement plate tower, or any other "contacting device" such
as a moving bed or venturi scrubber. The vapor-liquid equilibrium
relations and the process chemistry involved for each specific S02
absorption process are important factors which directly affect the
design and hence the cost of the absorption system for that process.
Capital costs for both a slurry-based and a solution-based absorption
system are provided in Appendix B.
The SO,, regeneration section, or the disposal section if the process
is non-regenerable, is also uniquely identified with a specific control
process. The capital and operating costs for these two sections, SO-
absorption and regeneration, together reflect the economic differences
between the basic control processes themselves. These costs have there-
fore been combined to present a single cost function, and are presented
under the appropriate SO™ control process section as a family of para-
metric curves relating gas flow, S02 concentration, and cost.
10
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2.6 AUXILIARY S02 UTILIZATION PLANTS
The regenerable S02 control processes, with the exception of the
citrate process which directly produces elemental sulfur, provide a
concentrated stream of sulfur dioxide gas. This stream can be utilized
in a number of ways:
1) as feed to a sulfuric acid plant
2) as feed to an elemental sulfur plant
3) for conversion to liquid S02-
Sulfuric acid plants are usually designed for an S02 concentration
of 4 to 10 percent, with any higher concentrations being diluted with
additional air. The economics of elemental sulfur plants favor concen-
trated S02 gas feed streams, while liquid S02 plants must include suit-
ably sized predryers for those S02 streams recovered by steam stripping
and condensation. Thus, there is a tendency for the control processes
under equal conditions to be preferentially matched with a particular
S02 utilization process. For example, the regeneration process of the
magnesium oxide system produces an S0« stream of around 15 percent
concentration. An auxiliary sulfuric acid plant would offer the most
straightforward utilization avenue. On the other hand, the Wellman-Lord
process readily yields an S02 product stream with a concentration level
up to 80 percent, a level attractive for the production of elemental
sulfur or liquid S02. However, under certain conditions, the selection
of an alternative utilization operation may be well justified in spite
of the economic penalties which may be involved.
Capital and operating costs for sulfuric acid, elemental sulfur
plants and SO- liquefaction plants have been developed; and appropriate
cost curves are presented in Appendices C, D and E, respectively.
2.7 CAPITAL AND OPERATING COST ESTIMATES
Capital cost estimates have been developed from basic published
data, but considerable adjustment, together with appropriate escalation,
has been necessary to provide a degree of compatibility and relationship
to smelter operations.
As noted in Section 2.2, smelter effluent gas received at the
boundary limits of the control process has been standardized at 600°F
temperature with a maximum particulate loading of approximately 0.1
gr/SCF and an SO., concentration of 0.03 percent. Capital and operating
costs of both the gas conditioning and S02 absorption sections have been
related only to the gas flow rate, although in the latter section, the
chemistry and kinetics of the specific process itself will also influence
11
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the costs. In actual practice, the capital costs of the SC^ absorption
system would be expected to increase somewhat with increasing SO,,
concentrations, as well as with increasing gas volumes, but since the
capital estimates for- this study are based largely on generalized and
adjusted data, the contribution from S0« concentration has been ignored.
The capital cost of the blowers or fans to move the gas through both the
gas cleaning and S0~ absorption sections to the stack has been allocated
entirely to the S02 absorption section although in practice they may be
located in front of the gas conditioning section. However, operating
power costs for these blowers or fans have been allocated to both the
conditioning system and the absorption section based on expected pressure
drops in each section. Variable operating costs have been based on
determined unit usages related to either the gas flows or the SO- rate
(in Ibs/hour) as appropriate.
In both the regeneration and utilization sections, both capital and
variable operating costs have been related only to the hourly SO^ rate.
Operating labor has been allocated separately to each of the functional
sections with incremental man-hours being assigned arbitrarily at
75,000 SCFM and 10,000 Ib/hr S02 rate. Maintenance costs and allow-
ances for local taxes and insurance have been taken as percentages,
appropriate for each specific process, of the total capital investment
for that process.
In providing for the effect of depreciation and other capital
related costs, the commonly accepted approach of using an amortization
factor ("Capital Recovery Factor") or a fixed percentage of the capital
investment based on expected life has not been adopted. Corporate
income tax rates exercise a significant influence on the company's
effective annualized cost. In addition, differences between capital
cost-intensive and operating cost-intensive control systems are also
highlighted by consideration of the tax rate. At the present tax rate
of 48 percent, the net annualized cost of an installed S02 control
system is provided by the relationship:
NET ANNUALIZED COST = (Capital Recovery Factor)(Total Capital Investment)
+ 0.52 (Annual Operating Costs) - 0.48 (Annual Depreciation Rate).
For the purposes of this study the Annual Depreciation Rate has been
taken on the basis of straight line depreciation and not on the more
12
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usual industrial practice of sum-of-the-year's-digits depreciation. It
should be noted that if evaluations are desired excluding the effect of
income tax rates, the annualized cost becomes:
ANNUALIZED COST = (Capital Recovery Factor)(Total Capital Investment)
+ (Annual Operating Costs).
This information is readily available from the control cost tables
provided for each smelter.
The Total Capital Investment data developed for each process include
a 35-percent allowance to cover engineering, contractor's fees and
contingency contributions. Cost data for years prior to mid-1974 have
been escalated to mid-1973 at an annual rate of 5 percent. A rate of 15
percent has been used to escalate this date to mid-1974.
In the preparation of the final capital and operating cost curves
for each SO,, process, as noted in Section 2.5, the SO,, absorption and
the S0_ regeneration (or disposal) sections have been combined to provide
a family of curves relating gas volume and S02 concentration with cost
for each process. In addition, the separate cost curves for each of the
process sections—i.e., gas cleaning and conditioning, S02 absorption,
S02 regeneration or disposal, and S02 utilization—have been provided to
allow flexibility in defining an overall S02 control system to satisfy
different conditions.
It should be noted that the capital costs determined above are for
new installations and do not include any provision for retrofitting to
i an existing smelter. Also, no specific provision has been made for
reheat of the scrubbed gas discharged from the Sp2 absorber. If this
effluent cannot be discharged and effectively distributed using the
discharge fans and the high stacks commonly used in smelters, some
degree of reheating may be necessary. In view of the current energy
shortage, the use of direct-fired oil or gas burners seems to be precluded,
with emphasis being directed to some appropriate heat recovery system to
provide the necessary reheat. The cost and operating uncertainties
associated with this approach have suggested that this factor be omitted
from both the general cost curves for the S02 control processes and in
the actual matching of these S02 control processes with specific U.S.
copper smelters.
13
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To provide direct cost comparisons between the different SO,,
control processes, a total system capital and annual operating cost
breakdown has been prepared in tabular form for each control process
treating a range of gas flows containing 1.0 percent SO,,. In the case
of the regenerable processes, the most common S02 utilization section as
determined by past practice has been selected to complete the total
system. Also included are unit capital and operating costs on the basis
of gas flow (SCFM) and annual tons of SO^ removed.
2.8 PRODUCT DISPOSAL OR MARKETING CONSIDERATIONS
All S0~ control processes produce a product which includes the
sulfur values captured from the treated gas stream. This product may be
a water/solids slurry with the sulfur values held predominantely as
sulfates, sulfides and bisulfites as is the case with the "throwaway"
processes, or it may take the form of elemental sulfur, sulfuric acid or
liquefied or gaseous S0~ from the regenerable type processes. Whatever
the final product, it must be disposed of. If it cannot be used or sold
for other processing sequences, it must be discarded as a waste material—
usually at a cost which may include capital outlays as well as routine
operating expenses.
In the "throwaway" processes considered—limestone scrubbing and
the double-alkali processes—since there is no commerical use for the
sludge end product under present U.S. conditions, appropriate costs have
been provided for the disposal of the sludge and included in either or
both the capital and operating cost structures.
The sulfur products which may be produced by the regenerable processes
have commercial value and potentially they offer a source of income
which could offset the costs associated with operating the control
process itself. How much this contribution might be will vary with the
geographic location of the source, transportation facilities, seasonal
demands, size of the market and potential for growth, to name a few of
the considerations. In short, only an in-depth market survey spanning
the options offered by different sulfur end products from regenerable
S0_ control processes and the special conditions of the various smelter
regions can provide a meaningful response to the income contribution
question raised above. Scope limitations on this current study have
precluded this approach but a brief consideration of the impact of the
14
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byproduct credit for that product most commonly associated with each of
the regenerable control processes on the net total annualized cost has
been provided In Table 2.2.
In Section 12, those U.S. primary copper smelters presently
generating weak SC^ streams have been reviewed, and-these streams
characterized in terms of flow rate and S02 content. S02 control processes
have been appropriately matched with each of these smelters and capital
and operating costs developed from the earlier established cost data.
The cost structure of any control system, particularly the gas
conditioning and S02 absorption sections, installed on existing process
equipment in an on-line operating facility is strongly influenced by
such factors as the particular site conditions, the age of the plant and
equipment, the flue system and the physical layout of the equipment
itself. Generally, the sulfur handling section can be located in-
dependently of the existing equipment and cost is not adversely influenced.
The combined contribution of the above factors may increase the capital
cost of the conditioning and absorption installation by factors ranging
between 1.1 and 1.8 or even higher under certain circumstances. Retro-
fitting, and the associated cost penalties, are thus associated uniquely
with each specific plant. In developing the costs in Section 12, in-
formation on plant age, equipment and facility layout plots, and comments
from the plant operators themselves have been reviewed; and using general
engineering judgment, a "retrofit factor" or cost multiplier for new
plant installation costs has been defined for the gas conditioning and
S02 absorption sections for each smelter. Specific considerations are
discussed in Section 12 under each individual smelter.
Individual cost computation sheets for each control process and
each smelter have been provided in Appendix G. It is emphasized, however,
that these details have been provided essentially for comparative purposes
and additional detailed input is necessary to establish design needs and
firm engineering cost estimates.
15
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Table 2.2. EFFECT OF BYPRODUCT CREDIT ON NET TOTAL ANNUALIZED COSTS OF THE
REGENERABLE S02 CONTROL PROCESSES WHEN TREATING A GAS STREAM CONTAINING 1
% SO,
CONTROL PROCESS
Usual Byproduct
Av. Market Value
Annual Production at SCFM
70,000
100,000
200,000
300,000
Net Total Annualized
70,000
100,000
200,000
300,000
Production cost , $/ton
(Deb it) /Credit 70,000
$/ton 100,000
200,000
300,000
WELLMAN-LORD
H2S04
$40/ton
40,800 TPY
58,300
116,600
175,000
1,764,400
2,306,700
3,958,900
5,268,500
$43 ($3/ton)
$40 0
$34 $6/ton
$30 $10/ton
MAGNESIUM
OXIDE
H2S04
$40/ton
38,600
55,200
110,500
165,800
1,699,600
2,214,300
3,667,900
4,764,300
44 ($4/ton)
40 0
33 $7/ton
29 $ll/ton
AMMONIA
H2S04
$40/ton
40,800
58,300
116,600
175,000
1,831,800
2,357,800
3,701,100
4,796,200
45 ($5/ton)
40 0
32 $8/ton
27 $13/ton
CITRATE
Sulfur
$50/ton
13,700
19,600
39,200
58,900
1,359,900
1,731,200
2,764,000
3,575,600
99 ($49/ton)
88 ($38/ton)
70 ($20/ton)
61 ($ll/ton)
DMA
Liquid S02
$110/ton
27,400
39,200
78,500
117,800
2,115,800
2,785,100
4,742,600
6,429,000
77 $33/ton
71 $39/ton
60 $50/ton
55 $55/ton
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2.9 REFERENCES
1. Connor, J. M., "Economics of Sulfuric Acid Manufacture," Chemical
Engineering Progress, Vol.64, No.11, November 1968.
2. Semrau, Konrad T., "Feasibility Study of New Sulfur Oxide Control
Processes for Application to Smelters and Power Plants, Part II,
the Wellman-Lord S02 Recovery Process."
3. Fluor-Utah, "The Impact of Air Pollution Abatement on the Copper
Industry," PB-203-293 (1971).
17
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18
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3.0 SULFURIC ACID
3.1 PROCESS DESCRIPTION
The manufacture of sulfuric acid from smelter gas streams involves
essentially four steps: gas conditioning, drying, acid making, and acid
storage. Figure 3-1 shows a typical flowsheet.
1) Gas Conditioning
Effective removal of particulates and other impurities from the
smelter gases is essential for the operation of the acid plant, avoids
costly shutdowns and maintenance due to catalyst fouling and equipment
corrosion, and reduces the chances of acid contamination. Gas con-
ditioning also entails gas cooling.
Various equipment combinations for gas conditioning have been
proposed. However, for this study, a scrubbing tower, cooler, and electro-
static mist precipitator were used as the prime conditioning equipment.
This equipment not only removes the particulates and other impurities
but also cools the gas stream down to around 130°F (55°C).
2) Drying
The gases must be dried prior to conversion. However, some moisture
in the gas stream is helpful during the conversion stages. In sulfuric
acid plants utilizing gas streams with sulfur dioxide contents exceeding
3.5 to 4 percent, a drying tower using 93 percent sulfuric acid will
normally be capable of drying the gas stream. The removal procedure in
modern sulfuric acid plants is to dry the gases to moisture contents
below what is needed and then to add moisture prior to conversion. This
procedure results in better control of the moisture content of the gas.
For sulfuric acid plants utilizing gas streams with sulfur dioxide
contents of less than 3.5 to 4 percent, refrigeration may be required to
sufficiently dry the gases. The amount of drying needed is based on the
moisture-balance temperature, the temperature to which the gas stream
must be lowered to contain the proper moisture (see Section 3.2.3).
After refrigeration the gases may be further dried in a drying
tower, and then a carefully controlled amount of water is added in prior
to conversion.
19
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3) Acid Making
After leaving the drying section, gases then proceed to the acid
making section. Prior to the catalytic converter, the gas temperature
must be raised to 800-860°F (425-460°C). To accomplish this temperature
rise, the gas is passed through a series of heat exchangers utilizing
waste heat generated during the exothermic conversion of SO,, to SO..
In sulfuric acid plants processing streams containing 4 percent or
greater sulfur dioxide, autogenous operation occurs (the heat supplied
by the conversion of S0? to SO., is sufficient to keep the plant in heat
balance). For less than 4 percent SO-, heat must be supplied to raise
the temperature of the gases to the required conversion temperature.
This makeup heat is supplied by a continuously operating furnace. Large
amounts of energy (up to 24.6 million Btu per hour for a 0.5 percent SO-
stream at 100,000 SCFM) are required. Associated with this high energy
requirement is the high cost of the oil necessary for its production.
The gas stream is then introduced into the catalytic converter.
Normally 3- or 4-stage converters utilizing vanadium pentoxide catalysts
are used. The gas passes through the first stage of the converter where
the SO- is partially converted to SO-, with an associated temperature
rise, and then through a heat exchanger where it is cooled prior to
passing through the second catalyst stage. In single-absorption plants,
this process proceeds through the final converter stage where the gas is
then passed through a heat exchanger and then to the absorption tower.
In double-absorption systems, the gas stream is usually taken from the
converter after the second or third stage, passed through a heat exchanger
to be cooled, and then introduced into an absorption tower. The gas
minus part of the absorbed SO- is then reheated and reintroduced into
the converter where it follows the same path as in a single-absorption
plant.
In the absorption towers, the SO- is absorbed by strong sulfuric
acid and water is added to make the desired acid strength. From the
towers, the acid passes through an acid cooler and then to acid storage
tanks. The unabsorbed SO. and the remaining SO. In the gas stream are
vented to the atmosphere.
20
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REFRIGERATION BOOSTER
r
ABSORBER
777ff
77//7
////y
•
HEAT
EXCHA
BANK
FURNACE
TO STACK
MGER
T
ABSORBER j
ACID
PUMP
TANK
DUALABSORBTION
SECTION
DILUTE FEED GAS
SULFURIC ACID PROCESS
ACID
' WATER I COOLING
FIGURE 3-1
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4) Acid Storage
Acid from the absorption towers is normally passed through acid
coolers and then to carbon-steel storage tanks. A 30-day storage capacity
is normally sufficient.
3.2 PROCESS AND OPERATING CONSIDERATIONS
Sulfuric acid plants are basically steady-state systems. Those
operating on smelter gas streams are subjected to transient conditions
that include drastic changes in both gas flows and sulfur dioxide con-
centrations. Other significant operating considerations based on gas-
stream parameters include the presence of highly active metallic and
volatile compounds, the sulfur dioxide content, the moisture content,
and the temperature.
3.2.1 Transient Operation
As mentioned above, sulfuric acid plants operating on smelter gases
are subjected to transient conditions as a normal operating condition.
Normally, sulfuric acid plants that are designed for steady-state con-
ditions, within restrictive margins, cannot tolerate these transient
conditions; and the plant either produces dilute acid or is shut down.
These conditions result in the release of large amounts of sulfur dioxide
to the atmosphere unless the smelter is shut down or an alternative
source of sulfur dioxide is used.
To avoid this situation, plants designed for transient operations
are "oversized" both in terms of gas handling capability and acid pro-
duction. This approach imposes economic penalties both in terms of
capital investment and operating costs. As noted in Section 3.1, acid
plants treating gas streams containing less than approximately 4.0
percent S0~ are not autogenous, and supplementary heat input to the gas
stream is necessary to sustain the S02 to SO, conversion. Variations in
gas flow and S02 content will require special control loops to balance
and control this auxiliary heat input step.
3.2.2 Metallic Particulates and Fumes
Although cleaning and conditioning of the gas stream prior to
entering an acid plant is extensive, submicron particulates and metallic
fumes are carried through to the converter. Such materials will contribute
to plugging, poisoning or partial deactivation of the catalyst, materials
corrosion, and/or acid contamination.
22
-------
In addition, some metallic particulates such as iron oxides may act
as catalysts that promote the conversion of S09 to S0~ prior to gas
<- J
conditioning. The SO., thus formed can be removed from the gas stream
during the gas conditioning stage with a subsequent reduction in sulfur
content of the gas stream entering the converter.
3.2.3 Sulfur Dioxide Concentration
The market for sulfuric acid is generally focused on acid strengths
of 93 percent and greater. The concentration of acid produced is a
function of both the S02 concentration and the moisture content of the
incoming gas stream. With SO^ concentrations below about 3.5 to 4.0
percent, the equilibrium moisture content is more than enough to produce
93 percent or stronger sulfuric acid, and special treatment is required
to remove the moisture. The direct effect of the S02 concentration on
the degree of cooling that is necessary to adequately dry the incoming
gas can be seen by the moisture balance equation:
H <_ (98-80m)
c
PSO
2
221.9 m
where H is the saturated moisture content corresponding to the moisture
balance temperature, m is the desired acid concentration, x is the
S02/S0., conversion ratio, and P_ is the partial pressure of the sulfur
dioxide. 2
As the sulfur dioxide concentration drops, the saturated moisture
content drops, and accordingly, the temperature to which the gas stream
must be cooled drops. Refrigeration must be used to cool the gas stream
to the moisture balance temperature. Refrigeration, systems are energy
intensive, and this pretreatment requirement for low SO,, gas streams
imposes significant penalties in capital and operating costs.
Low SO. gas streams also require the use of auxiliary heat to
insure proper operation of the sulfuric acid plant (see Section 3.1.3),
Associated with low S02 concentrations, high volumetric flows are en-
countered in smelters. To handle these flows, additional expense is
required for both larger air handling equipment and for increased energy
requirements associated with the larger equipment.
23
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A viable alternative to the direct use of low sulfur streams is
upgrading the streams prior to conversion. This can be accomplished by
effective draft control by redesign or cleaning of hoods and flues; by
gas recycling; and by reduction of air infiltration by filling crevices;
by shortening burner-to-refractory distances, and by tightening and
sealing flues and hoppers. The application of these methods has resulted
in a 70 percent reduction in air infiltration and a 32 percent increase
in the SO- content of the flue gas by one Japanese smelter. However,
the scope of this project precludes the examination of such alternatives.
3.2.4 Moisture Content
The water/SO- ratio is important because of the equimolar quanti-
ties of water and SO- that are needed to complete the reaction
More water can be tolerated for lower strength acids according to the
formula for the water balance temperature mentioned previously. In most
cases, a drying tower follows the refrigeration unit in low sulfur
streams to further dry the gases. After the drying tower, the correct
amount of water is usually injected into the gas stream prior to the
converter. This procedure allows for a more precise water/S02 ratio
adjustment to produce the desired acid concentration.
3.2.5 Temperature
The temperature of the gas stream is very important in many stages
of acid making. The gas stream leaves the reverberatory furnace at
about 1200°F (650° C) and is cooled in a waste heat boiler and a spray
tower or other gas conditioning device. Intermediate stages of cooling
using adiabatic direct cooling by evaporation or gas/liquid heat ex-
changers may also be needed. From the gas conditioning, the gas stream
at about 130°F (55 °C) must be dried, and in the case of low sulfur
streams (<4 percent), must be refrigerated to reach the moisture-balance
temperature.
After cooling to remove the water, the gas must be heated to 800-
850°F (425-455 °C) prior to entering the converter. This is accomplished
by utilization of the heat generated from the conversion of SO- to SO-
24
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for autogenous plants (those using streams of approximately 4 percent
S02 or greater), while supplementary heat must be applied for lower SCK
streams (<4 percent SCL).
The cooling and heating must be accomplished rapidly, especially
prior to the final gas cleaning, because dust particles in the gas
stream at temperatures in the range of 830-1100°F (445-595°C) may act as
catalysts for the formation of S0» with the subsequent sulfur yield loss
of the gas stream. Therefore, short contact periods between the gas and
dust at these temperatures are essential for low S0_ formation.
3.3 PROCESS DEVELOPMENT STATUS
Sulfuric acid plants have been applied to high S02 smelter gases
for many years. The success of their application to low SO,, gas streams
(<3.5 to 4 percent) is questionable. In 1971, a sulfuric acid plant was
applied to the treatment of the exit gas from a green-charged rever-
beratory furnace at Onahama Copper Smelter near Onahoma, Japan. Prior
to this application, the S0~ content of the gases was upgraded from
about 1.5 percent to about 2.2 percent.
The article which described this application also stated the
following:
When producing concentrated acid from weak gases by a contact
process, the practical lower limit (not the economical limit)
of S02 concentration is 2.0 percent judging from such factors
as the temperature limit to maintain the water balance,
supplemental heat to the conversion system, etc.
Judging from this statement and from conversations with sulfuric acid
plant designers, application of this process to weak SO^ streams from
smelters without prior enrichment processes would not seem to be reasonable.
3.4 PROCESS DESULFURIZATION EFFICIENCY
The theoretical equilibrium conversion of S02 to S0_ for acid
making can exceed 99 percent and is related.to both the temperature of
conversion and the SO- content. Generally, the lower the temperature
and the lower the S0~ content, the higher the conversion efficiency.
Increasing the oxygen concentration and/or pressure also increases the
conversion efficiency, but the relative gain is small and much higher
2
costs are realized.
25
-------
Favorable low temperatures have correspondingly unfavorable reaction
rates. At about 750°F (400°C), the reaction rate is nil with a vanadium
catalyst, and fairly slow at about 840°F (450°C). Consequently, the
gas is heated to around 800-860°F (430-460°C) prior to entry into the
converter. The lower temperatures cause lower conversion efficiencies
which are compensated for by the use of a multiple-stage converter
(usually three or four stages).
Theoretical conversion efficiencies are not obtained, but the use
of the multiple-stage converters can account for the conversion of 97 to
98 percent of the S02 to S03 which is then absorbed into sulfuric acid.
The use of double absorption can result in the conversion of at least
3
99.5 percent of the original SO™ concentration.
These conversion efficiencies occur only at peak conditions and
will be reduced considerably by both fluctuations in flows and S0?
content and by an aging catalyst.
3.5 PROCESS ENERGY REQUIREMENTS
The energy requirements of the sulfuric acid plants vary according
to both the gas flow and the sulfur dioxide concentrations. At low
sulfur dioxide concentrations (<4.0 percent) the energy costs are ex-
tremely high. For example, at 150,000 SCFM the energy costs are $750,000
per year for a 4 percent S02 stream; while for a 1 percent stream at
150,000 SCFM, the costs are over $4 million, an increase of approximately
550 percent.
The increase in energy costs is primarily due to the electrical
power requirements of the refrigeration unit needed for the 1 percent
stream, 83 percent of the total increase. The makeup heat required by
the 1 percent stream accounts for the remaining 17 percent.
3.6 PROCESS COSTS
Data on the costs of sulfuric acid plants operating on low sulfur
dioxide gas streams are practically non-existent because of the lack of
operating plants. This lack of data has led to the derivation of cost
estimates for capital and operating costs (Figures 3-2 and 3-3).
The total capital costs for acid plants are based on the sum of the
costs of acid making and gas cleaning, refigerating, and preheating.
26
-------
Acid making and gas cleaning — The first step in the development
of capital costs for sulfuric acid plants operating on low sulfur dioxide
gas streams was the generation of a cost curve for an 8 percent plant.
This curve was based on cost data collected from many sources and updated
to mid-1974 costs.3'4*5
The cost was divided into acid making and gas cleaning costs (Figures
3-4 and 3-5). For the acid making section, it was assumed that one half
of the section was flow dependent and one half was sulfur dependent.
This assumption, along with an assumed size factor of 0.65, resulted in
the following equation as an expression of the cost of the acid reaction;
where V was the volumetric flow rate (SCFM), C was the cost ($10 ),
subscript x was the desired SO- percentage, and subscript 8 was at 8
percent S02.
For gas cleaning, it was assumed that all the gas cleaning cost is
volume dependent and that a size factor of 0.68 applied. These assump-
tions resulted in the following equation.
V \ 0.68
x ] (cQ) - c
VQ / VW8' x
O '
Refrigerating — For the regrigeration costs (Figure 3-6), the
following assumptions were made: no regrigeration is needed for a 4
percent or greater S02 concentration, and the gases must be cooled from
130°F (55°C) to the moisture balance temperature. The cost of refrig-
eration was obtained from the open literature and a vendor. '
Preheating — The cost of preheating was assumed to be 15 percent
of the acid making cost for plants operating on less than 4 percent S02
gas streams (Figure 3-7). Preheating was considered unnecessary for
those plants operating on 4 percent or better SO- streams.
27
-------
The annual direct operating costs are shown in Figure 3-3. These
costs are made up of the operating costs for the various segments of the
process (Figures 3-8 through 3-11), plus the labor, maintenance, insurance
and taxes. Typical operating costs for various flow rates at 1 percent
SO- are shown in Table 3.1.
3.7 ADVANTAGES AND DISADVANTAGES
The chief advantages for sulfuric acid plants for control of SO-
from smelters are:
1) Relatively high S02 control efficiency.
2) Marketable product under most circumstances.
The chief disadvantages of sulfur acid plants for control of weak
S02 streams from smelters are:
1) Lack of operational plants to properly evaluate the actual
operating effectiveness on low sulfur dioxide streams.
Only one plant has been applied to this type of situation
and it is no longer operating.
2) High capital and operating costs.
3) The catalyst is susceptible to deactivation and fouling from
trace metals that are associated with smelter offgases.
4) Process is not autogenous at low S02 concentrations, i.e.,
makeup heat in large amounts and at high costs is needed.
5) At low S02 concentrations, refrigeration, with its
associated high energy requirements and costs, is required
to dry the gases prior to conversion of SO- to SO-.
6) Fluctuating flow rates and SO- concentrations and
deteriorating catalysts cause reductions in SO- removal
efficiency.
7) Cooling water demands are high.
3.8 REFERENCES
1. Niimura, M., T. Konada, and R. Kojima, "Sulfur Recovery from Green-
Charged Reverberatory Offgases at Onahama Copper Smelter," Paper
No. A73-47, presented at Metallurgical Society of AIME Mtg., 1973.
2. Duecker, W.W., and J.R. West (eds), The Manufacture of Sulfuric
Acid, Am. Chem. Soc. Mono. Series, Reinhold Publishing Corp, New
York, 1959.
28
-------
3. Chemical Construction Corporation," Engineering Analysis of Emissions
Control Technology for Sulfuric Acid Manufacturing Processes,"
NAPCA Contract No. CPA 22-69-81, March 1970.
4. Connor, J.M., "Economics of Sulfuric Acid Manufacture," in Chem.
Eng. Prog., Vol. 64, No. 11, Nov. 1968, pp. 59-65.
5. Fluor Utah Engineers and Constructors, Inc., "The Impact of Air
Pollution Abatement in the Copper Industry," published by Kennecott
Copper Corporation, New York, NTIS PB-208-293, April 20, 1971.
6. Guthrie, K.M., Process Plant Estimating Evaluation and Control,
Craftsman Book Company of Am., Solana Beach, Calif. 1974.
7. Robinson, R., York Industrial, personal communication, November,
1974.
29
-------
DILUTE FEED GAS
SULFURIC ACID PROCESS
TOTAL CAPITAL INVESTMENT COSTS
SO2 REMOVAL EFFICIENCY 96%
0.5% SO,
10.0 Million
SO,
4% SO,
8% SO,
NOTE: GAS COOLING & CONDITIONING
INCLUDED
Costs: Mid
1000
10,000
SO2 RATE IN INLET GAS (Ibs/hr)
30
FIGURE 3
-------
DILUTE FEED GAS
SULFURIC ACID PROCESS
TOTAL ANNUAL DIRECT OPERATING COSTS
SO2 REMOVAL EFFICIENCY 98%
0.5%
1.0%
1.50/
2.0%
$10.0 MILLION
$1.0 MILLION
NOTE: GAS COOLING & CONDITIONING
INCLUDED
Costs: Mid 1974
10,000
100,000
S02 RATE IN INLET GAS (Ibs/hr)
H
FIGURE 3-3
-------
DILUTE FEED GAS
SULFURIC ACID PROCESS
TOTAL CAPITAL INVESTMENT COSTS
ACID PLANT SECTION
C/5
8
H
Z
LLJ
2
I-
co
Q.
O
_l
<
o
S1.0 MILLION
Costs: Mid 19'
10,000
SULFUR DIOXIDE RATE (Ibs/hr)
FIGURE'
-------
DILUTE FEED GAS
SULFURIC ACID PROCESS
TOTAL CAPITAL INVESTMENT COSTS
GAS CLEANING SECTION
0.5%
$1.0 MILLION
10,000
SULFUR DIOXIDE RATE (Ibs/hr)
FIGURE 3-5
-------
DILUTE FEED GAS
SULFURIC ACID PROCESS
TOTAL CAPITAL INVESTMENT COSTS
REFRIGERATION SECTION
CO
in
O
o
UJ
CO
LU
<
H
Q.
O
$1.0 MILLION
10,000
SULFUR DIOXIDE RATE (Ibs/hr)
FIGUR
-------
DILUTE FEED GAS
SULFURIC ACID PROCESS
TOTAL CAPITAL INVESTMENT COSFS
PREHEATING SECTION
0.5%
$1.0 MILLION
10,000
SULFUR DIOXIDE RATE (Ibs/hr)
35
FIGURE 3-7
-------
DILUTE FEED GAS
SULFURIC ACID PROCESS
TOTAL ANNUAL DIRECT OPERATING COSTS
ACID SECTION
$1.0 MILLION
CO
CO
O
o
e?
tr
LU
o.
o
I-
o
LU
QC
0.5%
8.0%
10,000
SULFUR DIOXIDE RATE (Ibs/hr)
FIGURES
-------
DILUTE FEED GAS
SULFURIC ACID PROCESS
TOTAL ANNUAL DIRECT OPERATING COSTS
GAS CLEANING SECTION
0.5%
$1.0 MILL ION
4.0%
8.0%
10,000
SULFUR DIOXIDE RATE (Ibs/hr)
37
FIGURE 3-9
-------
DILUTE FEED GAS
SULFURIC ACID PROCESS
TOTAL ANNUAL DIRECT OPERATING COSTS
REFRIGERATION SECTION
to
co
CO
8
oc
LLJ
D-
O
U
LU
DC
2.C
$1.0 MILLION
10,000
SULFUR DIOXIDE RATE (Ibs/hr)
FIGURE!
-------
DILUTE FEED GAS
SULFURIC ACID PROCESS
TOTAL ANNUAL DIRECT OPERATING COSTS
PREHEATING SECTION
0.5%
S1.0 MILLION
10,000
SULFUR DIOXIDE RATE (Ibs/hr)
39
FIGURE 3-11
-------
Table 3.1 SULFURIC ACID
TOTAL SYSTEM CAPITAL AND TOTAL ANNUAL COSTS
TOTAL CAPITAL INVESTMENT
_ ;
Annual Cost
A. Direct Operating
1. Acid Section
2. Gas Cleaning
3. Refrigeration
4. Maintenance
5 . Labor
6. Taxes & Insurance
TOTAL ANNUAL DIRECT
OPERATING COST
B. Annual Capitalized
Cost
C. Net Total
Annualized Cost*
$
$/SCFM
$/ Annual ton
SO 9 Removed
$/yr/ton
S0? Removed
$/yr/ton
S02 Removed
70,000
9,000,000
129
317
570,000
120,000
1,300,000
360,000
200,000
230,000
2,780,000
98
1,180,000
2,340,000
83
Gas Flow Rate
100,000
11,300,000
113
279
830,000
160,000
1,850,000
450,000
200,000
290,000
3,780,000
93
1,490,000
3,090,000
76
- SCFM (§1% S02
200,000
17,800,000
89
220
1,650,000
300,000
3,750,000
700,000
200,000
450,000
7,050,000
87
2,340,000
5,440,000
67
300,000
23,300,000
78
192
2,500,000
450,000
5,550,000
920,000
200,000
580,000
10,200,000
84
3,060,000
7,620,000
63
*Rase.d otv CoTporate Ta-x. /Rate of
-------
4.0 LIME/LIMESTONE SCRUBBING
4.1 PROCESS DESCRIPTION
The removal of S02 from flue gases by calcium ion capture has had
extensive application in the power boiler field by wet scrubbing of the
gases with lime or limestone slurries.
Figure 4-1 shows a basic schematic of the process. The principal
Process steps are:
1) SO,-, Absorption — The cleaned smelter gas is prepared for
scrubbing by a conditioning tower which cools and humidifies the gases
by means of water sprays. The conditioned gas enters a multi-staged
scrubber which can be of the spray tower, turbulent contactor, or
venturi variety. The gas is contacted countercurrently by a lime or
limestone slurry prepared from milled and sized material.
2) Demisting — The gases which have been freed of SOp by ab-
sorption in the limestone slurry are trapped in a chevron mist eliminator
which is washed with clear water to prevent the escape of acidic droplets
into the atmosphere.
3) Liquor Loop Operation — The pregnant slurry is cycled to a
•Lagoon for settling and dewatering, returning the water to the absorber.
4) Limestone Handling — This system consists of field storage
and transfer of mined limestone to a milling and sizing plant for
alurrying and makeup in the absorber circulating loop.
4«2 PROCESS AND OPERATING CONSIDERATIONS
The chemistry of the limestone/S02 system is quite complicated. As
many as 28 chemical equations have been postulated by some authorities
to characterize the reactions involved. A simpler mechanism is suggested
by the following forms:2
S02 + H20 + H+ + HSO~
H+ + CaCO •*• Ca*4" + HCO~
Ca + HSO~ + 1/2H20 -»• CaSCy 1/2H20
H+ + HCO~
41
-------
In addition to the calcium sulfite formed, oxidation will also
produce the sulfate:
CaS03 + 1/2 02 -*• CaS04.
Solids deposition and scaling are persistent problems in this
process. The use of high solids concentrations (up to 15 percent)
favors the limestone utilization but results in erosion and increased
3
solids deposition. Scaling due to salt desupersaturation is particularly
objectionable when it occurs in the scrubbing or demisting equipment.
By seeding the slurry in a hold tank with solid calcium sulfate, the
desupersaturation can be controlled at a point not in the scrubbing
4
circuit. Scaling can also be reduced by using a high liquid-to-gas
ratio which lowers the concentration of sulfite and sulfate formed per
scrubbing cycle. The L/G needs to be much higher for a high SCL con-
centration as found in smelter gases than has been found adequate in
power plant practice. S09 removal efficiency is favored by high L/G and
2
by low inlet SCL concentration. Removal efficiency will decrease as
SO- concentration increases in the gas feed as encountered in smelter
operations.
Tests on a reverberatory gas stream at the McGill pilot unit were
made at varying S09 concentrations. At constant L/G, S0~ removal at 0.4
percent S0» averaged 87 percent; at 1.6 percent S09, average removal
2
efficiency was 72 percent.
The partial pressure of SO^ over the solution must be kept low for
maximum absorption. Equilibrium data are not available for the complex
system involved in lime/limestone scrubbing. In the case of the SO /
Ca2SO./ CaHSO- system, S02 partial pressure above the solution at 5 pH
is about 1 mm_Hg; at 7 pH it is practically zero. On the other hand,
the solubility of CaSO» decreases as pH increases which would cause
supersaturation and solids scaling at the point of limestone addition
where pH is high. Solubility of CaS03 is about 50 ppm at pH 6 and rises
to 2000 ppm at pH 4. For this reason, limestone addition and return of
lagoon overflow is made at a holding tank, and not to the scrubber
2
circulating loop. Some investigators find that addition of a weak
-------
S02
ABSORBER
so2
ABSORBER
I
I
GAS
CONDITIONING
SECTION
I
INLET
€)
I
WATER
TO STACK
ABSORBER
CIRCULATION
TANK
*-QJ
BALL MILL
UMESTONE
STORAGE SILO
SLURRY
TANK
TO SETTLING AND DISPOSAL
POND
RETURN
WATER FROM
SETTLING POND
LIMESTONE SCRUBBING PROCESS
FIGURE 4-1
-------
organic acid (e.g., benzoic) will increase the solubility of CaCO_
J 5
without volatilizing SCL from solution, thus increasing S02 absorption.
The sludges formed and needing disposal are largely CaSO™ with some
CaSO, plus salts of other cations found in the limestones. These are
Na, K, Mg, etc., salts. Since these are highly soluble they tend to
build up in the recirculated lagoon or thickener overflow. There is,
however, sufficient bleed as occluded liquid in the settled solids to
prevent saturation in the scrubber circuits. A disadvantage in using a
lagoon for solids separation in average rainfall areas is the problem of
excessive water load in the recycled supernatant liquid.
4.3 PROCESS DEVELOPMENT STATUS
In 1930 the London Power Company first used alkaline scrubbing of
flue gases by passing limey Thames river water once through the scrubber.
This circuit was later closed to avoid contamination of the river.
Gases from ore roasting plants were scrubbed in the 1960's in Japan and
the USSR. The Tennessee Valley Authority did pilot work in the United
States during the 1950's. Several power plant applications were installed
in the 1960fs.7
In the United States, 21 lime/limestone scrubbing installations
have been built or planned between 1968 and 1977. These are for power
utility plants totalling 11,482 MW generating capacity. In Japan,
there were nine installations built or planned between 1964 and 1974
6 8
with a total flue gas flow of 1.2 x 10 SCFM.
Improvements to the lime/limestone slurry process are being made
continuously. Prevention of corrosion calls for stainless steels where
metal contacts acidic solutions. Erosion is reduced by elastomer
linings of pumps, piping, and surfaces subjected to heavy slurry im-
pingement. High temperature surfaces are protected with acid proof
gunite or glass-flake reinforced polyester resin linings.
Disposal of waste solids remains a serious problem. The smelter
industry, which emits 2 x 10 tons of sulfur annually, could produce 25
x 10 tons of sludge if this process were universally applied in the
United States. This staggering statistic has stimulated study of possible
disposal such as dumping into abandoned mines or quarry pits, chemical
44
-------
fixation to improve bearing strength for landfill, or as a base for hard
surfacing and conversion of the sulfite sludge to useable products such
as oxidation to gypsum. Efforts to optimize limestone utilization
include use of organic acids to aid dissolution and grinding to fine
size to increase "surface to volume" ratio. Much work is going into
scrubber design both as to basic type and optimal configurations.
Some of the problems and constraints which are anticipated in
applying lime/limestone scrubbing to weak smelter streams are expressed
Q
below by the A. D. Little organization. These problems will require
extensive pilot and demonstration work on actual smelter gases for this
resolution:
1) Need for multiple staging will increase scrubber cost per
CFM as compared with utility boiler applications.
2) Variability of SO™ feed concentrations from roasters and
reverberatories will not permit continuous high collection
efficiency.
3) Difficulty of passing on the added cost of operating a
scrubbing system.
4.4 PROCESS DESULFURIZATION EFFICIENCY
Effectiveness of S02 removal is improved by maximizing the SO^ gas
phase mass transfer rate at the liquid interface by proper choice of
scrubber type and by use of high liquid-to-gas ratio. The other con-
trolling criterion is to increase the surface availability of the lime
or limestone by reducing the particle size of feed solids and by using
high slurry recirculation rates, high solids contents, and large liquid
holdups.
Tests in power plant applications have reported SO^ efficiencies of
from 70 to 85 percent using limestone (Meramec and Will County). The
Phillips Station of Duquesne Power using a calcium hydroxide scrubbing
system of Japanese design has reported 80 to 90 percent removal.
2
Tests conducted by SCRA, Inc., on a reverberatory furnace gas at
the Kennecott McGill, Nevada smelter showed S02 removal efficiencies of
72 to 85 percent. Efficiency was highest at low SO,, inlet concentration
and high L/G ratios.
Because of the higher SCL content in smelter gases, a single stage
contactor would soon be loaded with sulfite with attendant scaling.
Multi-staging for smelter gas desulfurization is mandatory.
45
-------
Scrubbing at McGill used several stages of contacting. Gases are
first passed through a venturi scrubber and then through a 2-stage TCA
column. Slurry was recycled in both contactors. It was found that
recycling improved desulfurization effectiveness.
The general conclusion which developed from operations at McGill is
that with optimum conditions it should be possible to achieve 90 percent
S00 removal from a reverberatory gas stream containing 1 percent S09
tL £•
provided:
1) Use is made of a venturi followed by six stages of absorption.
2) High L/G ratio is maintained.
3) A high quality limestone ground to a moderate fineness is
used.
4) The inlet SCL concentration can be leveled, to avoid swings
and peaks.
Surging of gas concentrations in smelting operations is a serious
problem. In contrast to utility boiler operation which hold conditions
relatively steady, charging of a reverberatory furnace can cause peaking
to five times the daily average SCK inlet concentration. This cannot be
handled by the absorber unless uneconomically excess capacity is built
into it.
Maintenance problems and operating upsets would probably limit on-
stream availability of this equipment to 85 percent.
4.5 SULFITE OXIDATION
Oxidation of calcium sulfite to sulfate normally occurs in the
presence of oxygen in the smelter offgases. If sulfate is allowed to
build up in the scrubber circuit, there is danger of eventual desuper-
saturation and formation of adherent scale in the scrubber internals and
pumping elements. If seeding of the slurry with sulfate crystals is
allowed to occur in a holding tank with sufficient retention time,
desupersaturation at this point can be controlled. Because of its
crystal form, CaSO^ improves clarifier operation and solids settling in
ponds. The Howden-I.C.I. Process deliberately introduced an "oxidizer"
stage before the thickener to take advantage of the improved dewatering
characteristics of the sulfate crystals.2 In Japan, byproduct markets
46
-------
for gypsum have existed which encouraged development of processes for
oxidation of calcium sulfite wastes. In some cases, even if a saleable
byproduct was not produced, conversion of sulfite waste to sulfate was
found desirable where land use restrictions favored gypsum ponding with
10 percent moisture over sulfite storage at 60 percent moisture.
Efforts to reduce oxidation to sulfates have included sealing of
effluent tanks, use of oxidation inhibitors, and reducing liquid rate.
Tests indicate that higher rates of oxidation result at lower inlet
liquid pH (5.7-5.9), but these lower pH levels also reduce the amount of
lime needed per mole of SO^ absorbed and, therefore, reduce the cost of
absorbent and the amount of waste sludge.
There is little information on what effect smelter gases would have on
oxidation rates. There are no commercial installations of lime scrubbing
of smelter offgases. Since oxygen contents are greater in smelter gases
than in power plant flue gas, higher rates of oxidation might be expected.
It is also believed that trace metals in smelter streams may catalyze
oxidation reactions. Pilot plant work on simulated and on actual smelter
gases, however, does not demonstrate significant oxidation effects
2 12
attributable to trace metals. '
4.6 FEED GAS PRETREATMENT
It is expected that smelter gases, which would be treated by slurry
scrubbing for S02 removal, would be precleaned in an ESP. Statnick
reports actual tests on a copper smelter converter gas cleaned at 94 to
97 percent efficiency, indicating outlet particulate loadings of 0.04 to
0.08 gr/SCF. Gas pretreatment as part of the S02 capture process,
therefore, would be primarily for cooling and humidification at 130°F to
prevent evaporation and scaling in the scrubber. This is discussed in
Section 2.4. Since the Limestone Process is a throwaway type, exhaustive
particulate removal is not warranted.
4.7 PARTICULATE REMOVAL CAPABILITY
The particulate loading in a gas stream which has been pretreated
in an electrostatic precipitator with greater than 96 percent efficiency
would be generally less than 0.1-0.2 grains/SCF with a size distribution
predominantly under 4-5 microns. Gas cooling and conditioning using
47
-------
predominantly under 4-5 microns. Gas cooling and conditioning using
spray or impingement towers will provide some degree of additional
cleanup so that the gas stream actually introduced into the absorption
section of an S0~ control system will be a relatively clean gas, with
any entrained particulate matter being in the submicron range. A normal
pressure drop absorption column, whether it is a sieve tray column,
packed tower or a turbulent contact type (TCA) column, will provide
little in the way of removal of sub-micron particles.
In the Lime/Limestone Process, venturi and TCA-type columns are
commonly used in the S02 absorption circuit, and although the pressure
drops used in the columns in this service are relatively low, the extremely
high liquid/gas ratios (around 100) necessary with gas streams containing
1 percent or greater S02 concentrations provide conditions under which
some additional fine particulate removal might be expected.
4.8 PROCESS ENERGY REQUIREMENTS
Power requirements for the Lime/Limestone Process constitute
approximately 15-20 percent of the total annual direct operating costs
although they are related to the S02 concentration in the gas stream.
The limestone must be ground to approximately 70 percent -200 mesh, and
with stoichiometric rates of approximately 150 percent, richer S09 gas
streams directly increase the power requirements. As SO- concentration
increases, liquid/gas ratios must also increase, or alternatively,
multi-stage absorption must be used with increased power requirements
for circulation and slurry transfer.
4.9 RETROFITTING REQUIREMENTS
The Lime/Limestone Process is extremely demanding in its overall
requirements for land space. Although the sludge disposal pond can be
located several miles, if necessary, from the actual control site,
generally the Mmestone grinding and slurry preparation areas must be
reasonably close to the scrubber system to minimize slurry-handling
problems. Slurry pipelines should be laid out as straight as possible
to minimize "settling-out" points and areas of local erosion.
The gas cooling section and the absorber section, together with the
slurry recirculation tanks and associated recirculation piping for this
process, are expected to require a larger than normal area for similar
48
-------
equipment to facilitate the special maintenance problems usually
associated with slurry systems.
There will be some variation in space requirements depending on gas
volume and SO- content, particularly in the limestone stockpile area,
grinding, and storage of ground material; but typical space requirements
have been estimated as follows:
Gas Conditioning and SCL Absorption:
(including slurry recirculation) 15-20,000 sq. ft.
Limestone Storage and Grinding: 10-15,000 sq. ft.
Sludge Disposal Pond: 100-300 acres
(depending on depth
and planned life)
4.10 PROCESS COSTS
Not surprisingly considering the attention given to this S02
removal process over the last 6-7 years, there is a plethora of cost
estimates in the general literature. Unfortunately, in addition to the
frequently uncertain basis for many of these estimates, they also reveal
the early cost optimism associated with developing processes. As the
Lime/Limestone Process has passed through pilot plant evaluations,
small scale installations, and full-scale commercial plants, the cost
profile has climbed steeply.
In developing a general cost structure applicable to the specific
gas/S02 content of nonferrous smelter gases, the approach outlined in
Section 2 was followed, with costs for this particular process being
developed independently for the limestone preparation area, the sludge
handling and disposal system, and a 2-stage SO, absorption section. The
cost for a total control system includes the cost for a secondary level
("throwaway") gas conditioning system. Reported cost information * ' '
was reviewed, adjusted as necessary to meet the objectives of this study
and escalated to a mid-1974 base. In particular, disposal pond costs
were adjusted downwards to reflect that only minimum site clearing and
excavation would be expected in the areas in proximity to smelter sites.
Many smelters are served by nearby open mining operations and these
49
-------
areas, in at least some cases, could provide sludge ponding capability.
It was judged that with gas streams containing in excess of 0.5 percent
SCL, a double SCL absorption column would be necessary, and the costs
reflect this approach.
The capital costs for various S0~ rates have been combined with the
capital cost of the absorption section to provide the parametric capital
cost relationships detailed in Figure 4-2. Total annual direct operating
costs based on gas flow and S0~ concentration are provided in Figure 4-
3.
Table 4.2 provides capital and operating cost breakdowns for the
Lime/Limestone Process applied to a range of gas flows containing 1
percent SCL. Table 4.3 provides a total system cost including gas
conditioning on the same basis. Both tables allow direct comparison
with the other candidate processes and their total system costs under
the same base conditions.
4.11 ADVANTAGES AND DISADVANTAGES
Advantages of this process are:
1) The process has the advantage of several years of
development through to full scale commercial installations.
Operating problems, materials of construction, and process
control considerations have been thoroughly evaluated and
process performance, at least in the public utility area,
can be now projected with some assurance.
2) Raw material requirements are relatively high but limestone
is generally a readily available material at smelter sites.
3) Capital and operating costs appear to be attractive in
relation to other control system options.
The disadvantages of this process include:
1) Process development and application has been almost entirely
in the public utility area and the conditions here favor a
process like the lime/limestone system. Although the gas
flows are high, they are constant, the level of S09 is
uniform and the close control of combustion conditions
keeps oxygen levels in the flue gas to moderate values.
This situation does not exist in nonferrous smelter
operations.
2) The process generates large quantities of a sludge 6-7 lbs/lbSO_
with unfavorable physical characteristics. Settling is difficult
and the space requirements are generally awesome. There are
few smelter locations which could handle the volumes
of sludge to be expected from a smelter operation over the
long term without some difficulty.
50
-------
3) Close process control, particularly pH control, is essential
to prevent or minimize scaling in the absorber and under
the varying gas flows and SO- concentrations usually
encountered in smelter operations, there is doubt how
effectively adequate control could be maintained.
4) S02 absorption efficiency falls off markedly as the S02
level in the gas stream increases and to compensate, very
high liquid/gas ratios and/or multi-stage absorption columns
are necessary with concomitant increases in both capital and
operating costs.
4.12. REFERENCES
1. Raben, I. A., "Status of Technology of Commercially Offered Lime
and Limestone Flue Gas Desulfurization Systems,"Flue Gas Desulfurization
Symposium, 1973.
2. "The Removal of Sulfur Dioxide from Copper Reverberatory Furnace
Gas by Wet Limestone Scrubbing," Smelter Control Research Assoc.,
Inc.
3. Weir, Alexander, "Scrubbing Experiments at the Mohave Generating
Station," Flue Gas Desulfurization Symposium, 1973.
4. Epstein, M., C. C. Levio, C. H. Rowland, and S. C. Wang, "The
Test Results from EPA Lime/Limestone Scrubbing Test Facility,"
1973.
5. Hatfield, J. D., and J. M. Potts, "Use of Weak Acids to Improve
Sulfur Oxide Absorption by Limestone Slurries," AIChE Meeting,
1972.
6. Elder, H. W., and W. H. Thompson, "Removal of Sulfur Dioxide from
Stack Gases: Recent Developments in Limestone Wet Scrubbing
Technology," Transactions of the ASME, July 1973.
7. Slack, A. V.., H. L. Falkenberry, R. E. Harrington, "Sulfur Dioxide
Removal from Waste Gases," Journal of the American Pollution Control
Association, March 1972.
8. Ando, Jumpei, "Status of Japanese Flue Gas Desulfurization
Technology," Flue Gas Desulfurization Symposium, 1973.
9. Berkowitz, Joan B., A. D. Little, Inc., response to questionnaire
by J. R. Farmer, EPA, February 20, 1973.
10. Ando, Jumpei, "Utilization and Disposing of Sulfur Products from
Flue Gas Desulfurization Processes in Japan," Flue Gas Desulfurization
Symposium, 1973.
11. Epstein, M., L. Sybert, S. C. Wang, C. C. Leivo, and R. G. Rhudy,
"Limestone and Lime Test Results at the EPA Alkali Scrubbing Test
Facility at the Shawnee Power Plant," Symposium on Flue Gas
Desulfurization, November 1974.
51
-------
12. Campbell, I. E., "Status Report on Lime/Wet Limestone Scrubbing
to Control S0? in Stack Gases," E/MJ, December 1972.
13. Statnick, R. M., "Measurement of Sulfur Dioxide, Particulate and
Trace Elements in Copper Smelter Converter and Roaster/Reverberatory
Gas Streams," EPA, 1974.
14. E. L. Calvin, Catalytic Inc., "A Process Cost Estimate for
Limestone Slurry Scrubbing of Flue Gas," Contract 68-02-0241,
January, 1973.
15. G. G. McClamery and R. L. Torstrick, "Cost Comparisons of Flue
Gas Desulfurization Systems," paper presented at Flue Gas
Desulfurization Symposium, Atlanta, Georgia, November 4-7, 1974.
16. B. G. McKinney, A. F. Little, J. A. Hudson, "The TVA Widow's
Creek Limestone Scrubbing Facility," paper presented at the 1973
Flue Gas Desulfurization Symposium, New Orleans, La., May 14-17,
1973.
52
-------
Table 4.1. LIME/LIMESTONE PROCESS
UNIT USAGE AND COST DATA
A. Chemicals 6 Utilities
Basis
Unit Cost
Limestone
Power a) Scrubbing
b) SO- handling
Water
3 lbs/lbS02
6.4 KW/M SCFM
0.09 KWh/lbS02
1.6 gal/lbS02
$4/ton
$0.015/KWh
$0.10/Mgal
B. Operating Labor &
Maintenance
Labor
Maintenance
1% man/shift
<75,000 SCFM
<10,000 lb/lbS02
2 man/shift
>75,000 SCFM
>10,000 lb/lbS02
plus 1 man(days)
5.0% TCI/yr
$8/hr
C. Fixed Charges
15.7% TCI/yr
Based on Capital Recovery Factor using 10% interest over 15 year life plus
% taxes and insurance.
53
-------
LIMESTONE SCRUBBING
TOTAL CAPITAL INVESTMENT COSTS
S02 REMOVAL EFFICIENCY 90%
(2-STAGE SO2 ABSORPTION)
4.0%
10.0 Million
NOTE: GAS COOLING & CONDITIONING
NOT INCLUDED.
0.2% S02
1.0% SO
2
0.5% s:
(SINGLE SL
ABSORPTK
Costs: Mid
100,000 SCFM
GAS FLOW RATE-SCFM
FIGURE
-------
LIMESTONE SCRUBBING
TOTAL ANNUAL DIRECT OPERATING COSTS
S02 REMOVAL EFFICIENCY 90%
(2-STAGE SO2 ABSORPTION)
4.0% SO,
1.0 Million $
0.5% SO-
0.2% S0
(SINGLE STAGE
ABSORPTION
EFF. = 80%)
NOTE: GAS COOLING & CONDITIONING
NOT INCLUDED.
Costs: Mid 1974
100,000 SCFM
GAS FLOW RATE-SCFM
55
FIGURE 4-3
-------
LIMESTONE SCRUBBING
TOTAL CAPITAL INVESTMENT COSTS
LIMESTONE HANDLING & DISPOSAL
LIMESTON;
PROCESS
O
O
UJ
z
_J
<
OL.
O
1.0 Million
D ISPO;
ca
Costs: Midli
10,000 LBS/HR
SULFUR DIOXIDE RATE
IN ABSORPTION SYSTEM (Ibs/hr)
56
FIGURE
-------
CAPITAL AND TOTAL ANNUAL COSTS
-
TOTAL CAPITAL INVESTMENT
Annual Cost
A. Direct Operating
1. Limestone
2; Power
3. Water
4. Labor
5. Maintenance
6. Taxes & Insurance
TOTAL ANNUAL DIRECT
OPERATING COST
B. Annual Capitalized
Cost
C. Net Total
Annualized Cost*
$
$/SCFM
$ /Annual ton
S02 Removed
$/yr/ton
S02 Removed
$/yr/ton
Gi
70,000
3,150,000
45
121
312,000
125,000
8,300
121,800
157,500
78,800
804,200
31
414,200
- 731,600
28
is Flow Rate - J
100,000
4,000,000
40
108
446,800
178,600
11,800
156,600
200,000
100,000
1,093,900
29
526,000
966,800
26
3CFM @1% S02
200,000
6,400,000
32
86
893,600
357,300
23,700
156,600
320,000
160,000
1,911,200
26
841,600
1,630,600
22
300,000
8,400,000
28
75
1,340,400
535,800
35,700
156,600
420,000
210,000
2,698,500
24
1,104,600
2,239,000
20
Ui
*Based on Corporate Tax Rate of 48%
-------
Table 4.3. LIME/LIMESTONE PROCESS
TOTAL SYSTEM CAPITAL AND ANNUAL OPERATING COSTS
00
A. TOTAL CAPITAL INVESTMENT
1. Gas Conditioning
2. Limestone Scrubbing
TOTAL
B. TOTAL ANNUAL
OPERATING COST
1. Gas Conditioning
2. Limestone Scrubbing
TOTAL
C. NET TOTAL
ANNUALIZED COSTS*
1. Gas Conditioning
2. Limestone Scrubbing
TOTAL
$/Annual ton
SO - Removed
$/yr/ton
S02 Removed
^/vr/ton
Gas Flow Rate - SCFM @1% S02
70,000
1,210,000
3,150,000
4,360,000
167
211,000
804,200
1,015,300
39
230,200
731,600
961,800
37
100,000
1,500,000
4,000,000
5,500,000
148
287,000
1,093,900
1,380,900
37
298,500
966,800
1,265,300
34
200,000
2,250,000
6,400,000
8,650,000
116
465,200
1,911,200
2,376,400
32
465,800
1,630,600
2,096,400
28
300,000
2,900,000
8,400,000
11,300,000
101
635,700
2,698,500
3,334,200
30
619,200
2,239,000
2,858,200
26
-------
5.0 SODIUM SCRUBBING .- REGENERATIVE (WELLMAN-LORD)
5.1 PROCESS DESCRIPTION
The Wellman-Lord Process provides a method by which the S02 from a
weak stream can be absorbed by chemical reaction with an alkaline
scrubbing liquor. The S0~ is later desorbed by a heating process in •
which the S0? appears in concentrated form and the absorbing solution is
regenerated for recycling to the scrubber.
The strong S02 gas resulting from the above absorption and desorption
can be further processed along one of the following routes:
1) Dry and cool the gas to produce liquid S0~
2) Feed the wet S02 to a sulfuric acid plant
3) Catalytic reduction and conversion to elemental sulfur
(Glaus reaction).
Figure 5-1 shows a simplified schematic of the process which consists
of the following stages:
1) Absorption
2) Crystallization
3) Solids Separation
A) Recycling.
Absorption — After particulate removal and cooling, the gas stream
enters a two or three stage tower. This is either a tray or a packed
unit. The gases are scrubbed by an alkaline solution of Na2SO,.
Crystallization — The pregnant absorbent is pumped to an evaporator/
crystallizer for regeneration of the Na2S03.
Solids Separation — The crystal magma from the crystallizer is
filtered out of the mother liquor.
Recycling — The Na2S03 crystals are dissolved and recycled to the
absorber.
5.2 PROCESS AND OPERATING CONSIDERATIONS
The absorption process chemistry is similar to the sodium-based
double alkali system described elsewhere in this study. The scrubbing
solution consists of soluble NajSO^, NaHSO-j, and NajSO^. The S02
reacts with Na2S03 as follows:
S02 + Na2S03 + H20 •*• 2NaHS03.
59
-------
The rich bisulfite absorber bottom is pumped to an evaporator/
crystallizer while the gas stripped of S02 leaves the top of the absorber.
Regeneration is proprietary with Wellman-Lord. The heated solution in
the evaporator/crystallizer is stripped of SO™ as Na2SO™ is regenerated:
AH
2NaHS03 ->• Na2S03 4- + S02 + + H™0 t.
The wet SO™ goes to a partial condenser for removal of water which
is re-used. The Na™SO~ crystal magma is centrifuged or filtered, with
£• J
the solids going to a dissolving tank for recycling to the absorber.
The concentrated SO™ stream is then fed to a sulfur plant, to a
sulfuric acid plant, or processed as liquid SO,,.
The liquid/gas (L/G) ratio is kept low in the absorber. The high
concentration of salts and low water load has the effects of lowering
oxygen absorption, reducing heat demand on the evaporator/crystallizer,
and reducing the size of vessels, pumps, and piping. Because of this
low throughput, the absorbent fluid is recirculated at each stage of the
tower to assure wetting of the tray or packing. Since absorption of
oxygen is liquid-film controlling, the low L/G ratio minimizes oxidation.
NaHSO™ is more soluble than Na™SO™. By feeding a strongly concen-
trated solution of Na2SO™ to the absorber, scaling and salt precipitation
are not experienced since reaction with S02 produces the soluble NaHSO™.
This same solubility difference operates to advantage in the stripping
operation: as SO™ comes out of the reactants, the resulting Na™SO™
£» £* j
crystallizes out, thus driving the stripping reaction forward.
Waste streams from this process consist of the lime neutralization
product from the feed gas cooling and conditioning step plus a purge
stream from the absorber circuit to eliminate the oxidation product,
Na2SO,. Loss of sodium ion from this purge is made up by adding NaOH to
the scrubbing liquid. Fractional crystallization equipment can be used
to separate the Na™SO^ from the sulfite/bisulfite fractions in the purge
and thus reduce sodium losses.
60
-------
I
GAS
CONDITIONING
SECTION
INLET GAS
€)
1
1
\
)
V
1
<
<
_J
^r
S02
=T
A
•— •
F
BS
Si
7
1
»W
~i
ORBE
JRGE
ANK
WE
r*
:R
! — t
N__f
-Ct
_y
\_
'N f
f
r
721
TO SULFURIC
AGIO PLANT
SULFUR
LIQUID SO,
50 % NoOH
STRIPPER
STEAM
CONDENSATE
WELLMAN - LORD PROCESS
REGENERATIVE SODIUM SCRUBBING
FIGURE 5-1
-------
Other users of the high concentration salt scrubbing solutions have
noted resistance to pH change by the buffering effect of the strong
4
sodium sulfite/bisulfite solutions. S02 concentration variation of 2-
1/2 to 1 in a kiln offgas showed no change in S02 capture efficiency.
This property is important in smelter applications where swings in SC^
concentration can be expected.
5.3 PROCESS DEVELOPMENT STATUS
i.i
Of the twenty installations built or planned for use of the Wellman-
2
Lord Process, half are in Japan. Application has been made to tail
gases of sulfur plants, sulfuric acid plants, and utility boilers fueled
by coal or oil; but none to smelter gases. The earliest installation in
the United States was at Paulsboro, New Jersey treating a sulfuric acid
plant tail gas. Experience with this installation in 1970 pointed up
problems with tube plugging, material corrosion and erosion, absorber
inefficiency, and accumulation of particulates in the regeneration
system.
Materials of construction have been changed to 316SS in all corrosive
and erosive areas. Feed stream to the regeneration step is filtered to
remove particulates. Absorber configuration was modified to improve
liquid/gas contact. Subsequent installations, especially in Japan, have
reported excellent reliability.
A carefully planned installation is scheduled soon to operate on
the stack gases of a 115 MW utility boiler of the Northern Indiana
Public Service Company. The Wellman-Lord absorption/desorption plant
will be followed by an Allied Chemical elemental sulfur plant. This
demonstration plant should provide definite answers to questions of
operation and economics.
5.4 DESULFURIZATION EFFICIENCY
SO- absorption efficiency is controllable by adjusting process
parameters at the absorber. In the Davy Power Gas Process, the absorbing
solution is recirculated around each stage to permit a low total feed
rate but a higher gas contacting rate at each tray.
An installation at Chiba, Japan treats a flue gas from oil-fired
boilers with inlet S02 concentration of 1500 ppm and exit at 150 ppm or
a collection efficiency of 90 percent. The design guarantee for the
62
-------
NIPSCO installation is also at 90 percent S02 removal. The latter
involves a utility boiler firing coal at 3.5 percent sulfur and will
reduce a 2300 ppm S0? flue gas to 200 ppm.
~ *
According to FMC Corporation the buffered, high concentration
sulfite/bisulfite scrubbing solution used in their double alkali process
can achieve 99 percent S02 removal by raising pH as required.
Potter and Craig report S02 collection of 92vpercent in the
ord, Paulsboro, New Jersey facility of the Olin Corporation
cleaning the tail gas of a sulfuric acid plant. The stream of 45,000
SCFM contains 6000 ppm S02 and is reduced to 500 ppm before venting.
Application to an oil-fired boiler of the Japan Synthetic Rubber Company
achieved 90 percent S02 absorption.
Actual experience of 90 to 92 percent S02 recovery using the
ellman-Lord Process is well documented. It also appears feasible to
raise this figure to a higher level if required. The criterion established
iT» the SCRA study2 is at 97 percent S02 removal.
5'5 SULFITE OXIDATION
In addition to the absorption and desorption operations indicated
ve> this process does produce some non-regenerable and therefore
esirable oxidation products which need to be suppressed or eliminated.
°ntact of the scrubbing solution with oxygen in the flue gas will yield
Na2S°4 according to:
Na2S03 + 1/202 -> Na2S04.
Th
e Presence of S03 from flue combustion gases will also produce Na2SO,:
2Na2SO + SO + H20 + Na^O^ + 2NaHS03.
emPeratures encountered in the evaporator/crystallizer, auto oxidation
e bisulfite will produce non-regenerable sulfate and thiosulfate
Co*Pounds as follows:
6NaHS03
Ua NaS0 kas to b
2S04 kas to be Purged from the system with loss of Na ion
IcH ^
18 replaced by injection of NaOH or Na-CO. into the scrubbing
Circuit.
63
-------
Suppression of Na2SO, formation or, at least, its minimization is
under active development. By controlling the process parameters and by
taking advantage of some of the inherent process characteristics, oxidation
effects have been reduced. Some of these methods are:
1) Maintaining a concentrated sulfite solution in the
scrubbing medium to reduce the solubility of oxygen.
2) Throttling the circulation rate in the absorber to reduce
the liquid/gas ratio to a value just necessary for S0~
absorption but not so high as to encourage oxygen to
liquid transfer.
3) Removing Na^SO, by fractional crystallization and
reacting with lime to precipitate gypsum and recycle
the NaOH.
4) Using antioxldants such as hydroquinone or hydrazine has
been suggested. Other proprietary chemicals have been developed
for this purpose.
Davy Power Gas claims that about 0.5 to 3 percent of the S09 in a
2
boiler flue gas entering the absorber will be oxidized. .Other estimates
3
are at 10 percent. Many weak smelter gases have higher oxygen content
than power plant flue gases due to stream dilution by infiltrating air
at ESP's, waste heat boilers, and leaky hoods and flues. These higher
oxygen concentrations will result in high S02 oxidation. The combined
effects of higher gas volumes and higher oxygen content adversely affect
capital and operating costs of gas scrubbing facilities.
5.6 FEED GAS PRETREATMENT
In this study it is assumed that the gas stream fed to the S0?
removal plant has been adequately cleansed of particulates by high
efficiency electrostatic precipitators or high energy venturi scrubbers.
It is also assumed that the high temperatures of the smelter gases have
been moderated by waste heat boilers to a level of about 600°F (316°C)
before entering the gas conditioning section of the plant.
As discussed in paragraph 2.4 and in Appendix A, gas conditioning
consists of adiabatically cooling the gas to about 130°F (54°C) in a
spray tower. This cooled, saturated gas enters the absorber in a state
which avoids evaporation of the scrubbing solution with consequent
deposition of solids. The spray tower circulating water is-neutralized
with lime and cooled by heat exchange. Any mist carryover is eliminated
by a mist precipitator.
-------
5.7 PARTICULATE REMOVAL CAPABILITY
As noted in Section 4.7, gas streams precleaned in dry electrostatic
precipitators and then conditioned via scrubbing have very low particu-
late mass loadings predominantly in the submicron range. Additional
treatment in mist precipitators prior to regenerable control systems
will tend to further reduce particulates. Thus the absorption section
of the Wellman-Lord Process will offer negligible capability in removing
additional particulates.
5.8 PROCESS ENERGY REQUIREEMNTS
With the absorption advantages of high concentration solution
scrubbing, liquid/gas ratios are significantly less than those required
for the Lime/Limestone Process and power requirements for absorbent
recirculation are correspondingly less. However, power requirements in
the forced-circulation evaporator/crystallizer section of the regeneration
area tend to compensate for this saving.
Steam requirements in the evaporator/crystallizer constitute a
significant proportion of the total energy needs of this process.
Single effect evaporation steam demands appear to be as high as 12
!b/ibso2 but the use of a double-effect unit will reduce steam con-
sumption by about 40 percent. Double effect evaporation has been
assumed for the smelter application although additional steam generation
facilities would still in all likelihood be required to support the
operation. Total energy requirements under the cost rates and usage
factors provided in Table 5.1 are approximately 40-45 percent of the
total annual direct operating costs for treating 1 percent S02 at the
rate of 100,000 SCFM.
5-9 RETROFITTING REQUIREMENTS
The regeneration section of this process can be located independently
°f the S02 absorption and gas conditioning sections, and this provides
some flexibility in retrofitting this process to existing plants.
Estimated space requirements are:
Gas Conditioning and S02 Absorption: 10-12,000 sq. ft.
S02 Regeneration Area: 8-10,000 sq. ft.
Auxiliary Plant (sulfuric acid
including storage): 25-30,000 sq. ft.
65
-------
5.10 PROCESS COSTS
,Capital costs for this process have „been developed using information
reported in cost estimates by the Tennessee Valley Authority and the
Lummus Company, discussions with Davy Power Gas Company, and general
cost estimation techniques. Following, the methodology adopted for this
study, the S02 absorption step was considered independently from the SQ?
regeneration and end use section, and the generalized S0'< absorption
cost curve for solution scrubbing provided'in Appendix B was taken as
directly applicable to the Wellman-Lord system. The SO., regeneration
section cost was developed against the background noted above with
appropriate cost escalation and allowance for indirect charges. The
cost curve versus S02 rate relationship is provided in Figure 5-4.
These two sets of data were then combined to provide the parametric
capital cost curves relating gas flow rate and S02 concentrations provided
in Figure 5-2.
Operating costs covering both the absorption section and the SO
regeneration section are based on the specific usage values and unit
costs provided in Table 5.1 and developed from the above sources. Total
annual direct operating cost curves for the range of S02 concentrations
and gas flows under consideration are provided in Figure 5-3.
5.11 ADVANTAGES AND DISADVANTAGES
The Wellman-Lord Process has wide applicability to utility and
industrial boiler flue gases and to acid plant or sulfur plant tail
gases. Although applicability to weak smelter gases has not been commerciallj
demonstrated, the following advantages of the process are noted:
1) S02 absorption capability is excellent and as a clear
solution scrubbing system, equipment maintenance and
operation poses no special problems.
2) The process generates a high purity stream of SO,—around
85 percent or higher if appropriate drying facilities are
added. This provides options among sulfuric acid,
elemental sulfur or liquid S0?.
3) The process has enjoyed considerable commercial application
and it appears that plants can be designed with a high
degree of confidence as indicated by the performance
guarantees being provided with the Wellman-Lord installation
for the Northern Indiana Public Service Company (NIPSCO)
on a 115 MW coal-fired utility boiler .
4) The process can handle wide swings in S02 concentration.
66
-------
The principal disadvantages of this system appear to be:
1) Oxidation in the absorbent results in purging and loss
of high cost sodium ions. The use of antioxidants appears
to be common in Japanese installations but this increases
the operating costs substantially. There are alternatives
such as fractional crystallization of the purge stream
to separate and regenerate the sodium sulfate, but this
approach is associated with both additional capital and
operating costs.
2) Total energy requirements of the overall process—i.e.,
gas conditioning, S02 absorption and regeneration, and the
necessary end-use plant—are appreciable and constitute a
major part (40-50 percent) of the total overall operating
costs.
5-12 REFERENCES
Schneider, R. T. and Earl, C, B., "Application of Wellman-Lord SO
Recovery Process to Stack Gas Desulfurization," Davy Powergas, 2
Inc., Lakeland, Florida. Paper presented at Flue Gas Desulfurization
Symposium, New Orleans, Louisiana, May 14-17, 1973.
2
Report to the U.S. Bureau of Mines by the Smelter Control Research
Association, March 1974, "Engineering Evaluation of Soluble Scrubbing
Systems."
3
LaMantia, C. R., R. R. Lunt, I. S. Shah, "Dual Alkali Process for
S02 Control." Paper 25 C, presented at the 66th Annual Meeting of
The American Institute of Chemical Engineers at Philadelphia, Pa.,
Nov. 11-15, 1973.
4
Brady, J. D., "FMC Corporation's Sulphite Absorption Process".
Presented at the Missouri Public Service Commission Conference,
September 1974.
5.
Potter, B. H. and T. L. Craig, "Commercial Experience with an S02
Recovery Process." Chemical Engineering Progress, August 1972.
Earl, C- B., Davy Power Gas Company, private communication.
McGlamery, G. G. and R. L. Torstrick, Tennessee Valley Authority.
Cost Comparisons of Flue Gas Desulfurization Systems." Paper
Presented at Flue Gas Desulfurization Symposium, Atlanta, Ga.,
N°v. 4-7, 1974.
67
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Table 5,1. SODIUM SCRUBBING REGENERATIVE - WELLMAN LORD PROCESS
UNIT USAGE AND COST DATA
A. Chemical & Utilities
Basis
Unit Cost
Sodium Carbonate: (Soda Ash)
Antioxidant:
Power: Absorption
SO? Regeneration
Water: a) Process Water
b) Cooling Water
Steam:
0.13 Ib/lb S02
0.002 lb/S02
4.0 KW/M SCFM
0.085 KWh/lbS02
0.2 gal/lbS02
1.2 gal/lbSO,
7.7 Ib/lbSO
-2
$52/ton
$2.00/lb
$0.015/KWh
$0.30/M gal
$0.19/M gal
$1.25/M Ib
B. Operating Labor &
Maintenance
Basis
Unit Cost
Labor: Absorption
S02 Regeneration
Maintenance:
Taxes & Insurance
% man/shift
<75,000 SCFM
1 man/shift
>75,000 SCFM
2% man/shift
<10,000 lb/lbS02
3Jg man/shift
>10,000 lb/lbS02
4.0% TCI/year
2,5% TCI/year
$8/hr
C. Fixed Charges 13.15% TCI/year
Based on Capital Recovery Factor, using 10% interest over 15 year life
68
-------
SODIUM SCRUBBING - REGENERABLE
WELLMAN - LORD PROCESS
TOTAL CAPITAL INVESTMENT COSTS
S02 REMOVAL EFFICIENCY 95%
4.0% SO,
$10.0 Million
0.5% S02
NOTE: GAS COOLING & CONDITIONING
NOT INCLUDED.
J-0 Million
Costs: Mid 1974
100,000 SCFM
GAS FLOW RATE-SCFM
69
FIGURE 5-2
-------
SODIUM SCRUBBING - REGENERABLE
WELLMAIM-LORD PROCESS
TOTAL ANNUAL DIRECT OPERATING COSTS
SO2 REMOVAL EFFICIENCY 95%
4.0% SO.
$10.0 Million
NOTE: GAS COOLING & CONDITIONING
NOT INCLUDED,
1.0 Million
Costs:
100,000 SCFM
GAS FLOW RATE-SCFM
70
-------
SODIUM SCRUBBING • REGENERABLE
WELLMAN-LORD PROCESS
TOTAL CAPITAL INVESTMENT COSTS
S02 REGENERATION SECTION
$1.0 Million
Costs: Mid 1974
10,000 LBS/HR
SULFUR DIOXIDE RATE
IN REGENERATION SECTION (tbs/hr)
71
FIGURE 5-4
-------
Table 5-2. SODIUM SCRUBBING REGENERATIVE - WELLMAN LORD PROCESS
CAPITAL AND TOTAL ANNUAL OPERATING COSTS
TOTAL CAPITAL INVESTMENT
Annual Cost
A. Direct Operating
1. Power
2. Water
3. Steam
4. Soda Ash
(Na2C03)
5. Antioxidant
6. Labor
7. Maintenance
8. Taxes & Insurance
TOTAL ANNUAL DIRECT
OPERATING COST
B. Annual Capitalized
Cost
C. Net Total
Annualized Cost*
*Based on Corporate Tax Rate
$
$/SCFM
$/ Annual ton
S02 Removed
•
$/yr/ton
S02 Removed
$/yr/ton
S02 Removed
of 48%
70,000
4,250,000
61
154
96,300
10,100
527,600
186,000
219,700
175,200
170,000
106,300
1,491,200
54
558,900
1,198,300
44
Gas Flow Rate
100,000
5,400,000
54
137
137,600
14,400
753,200
265,200
313,800
192,700
216,600
135,000
2,027,900
52
710,100
1,591,800
41
- SCFM @1% S02
200,000
8,600,000
43
110
275,100
28,900
1,507,300
531,300
627,700
262,800
344,000
215,000
3,792,000
48
1,130,900
2,827,500
36
300,000
11,210,000
37
95
412,800
43,200
2,259,600
795,600
941,400
262,800
448,000
280,000
5,443,000
46
1,472,800
3,945,000
33
-vl
N>
-------
Table 5.3. SODIUM SCRUBBING REGENERATIVE VTEL'LMAN-LORD PROCESS
TOTAL SYSTEM CAPITAL AND ANNUAL OPERATING COSTS
A. TOTAL CAPITAL INVESTMENT
1. Gas Conditioning
2. S02 Absorption &
Recovery
3. Sulfuric Acid Plant
TOTAL
B. TOTAL ANNUAL
OPERATING COSTS
1. Gas Conditioning
2. S02 Absorption &
Recovery
3. Sulfuric Acid Plant
TOTAL
C. NET TOTAL
ANNUALIZED COSTS*
1. Gas Conditioning
2. S02 Absorption &
Recovery
3. Sulfuric Acid Plant
TOTAL
$/ Annual ton
S02 Removed
$/yr/ton
SO 2 Removed
$/yr/ton
S02 Removed
70,000
1,800,000
4,250,000
1,420,000
7,470,000
280
273,100
1,481,200
199,500
1,963,800
74
321,100
1,198,300
245,000
1,764,400
66
Gas Flow Rate -
100,000
2,275,000
5,400,000
1,750,000
9,425,000
247
370,600
2,027,900
234,100
2,632,600
69
419,100
1,591,800
295,800
2,306,700
61
- SCFM @1% S02
200,000
3,550,000
8,600,000
2,675,000
14,825,000
195
613,500
3,792,000
271,200
4,776,700
63
672,200
2,827,500
459,200
3,958,90Q
52
300,000
4,650,000
11,200,000
3,450,000
19,300,000
169
843,300
5,443,400
462,000
6,748,700
59
740,000
3,945,000
583,500
5,268,500
46
*Based on Corporate Tax Rate of 48%
-------
7A
-------
6.0 DOUBLE-ALKALI SODIUM BASE (THROWAWAY PROCESS)
6.1 PROCESS DESCRIPTION
Although there is now a growing number of lime/limestone based
systems installed and in operation on utility power plants, the early
Problems of scaling, plugging, and equipment erosion gave impetus to the
investigation of soluble alkali scrubbing systems which will not precipitate
solids in the scrubbers and associated equipment. Regeneration of the
soluble scrubbing fluid takes place in separate reaction vessels under
controlled conditions. No slurries are in contact with the treated gas
stream as in the case of the lime/limestone systems.
Although double-alkali processes can be sodium, potassium, or
ammonium based, the interest and emphasis today is with the sodium and
ammonia systems. This section will focus on the sodium based system and
in particular the dilute absorbent alternative, although continuing
studies and the installation of commercial and pilot plant concentrated
systems suggest that this latter system has many advantages. Figure 6-1
Provides a simplified flowsheet representative of the dilute process.
The principal steps are:
1) Absorption — The cleaned, humidified flue or smelter gas
enters a mobile bed, valve tray or sieve tray scrubber,
where the gas is contacted by the recirculated absorbing
alkaline solution.
2) Regeneration — A portion of the pregnant absorbent is
pumped to a reaction tank where it is treated with lime
before flowing to a thickener or clarifier. The precipitated
solids are pumped to a rotary filter and discarded. The
overflow from the thickener is pumped to the softener.
3) Softening — Soda ash and frequently CO,, is added at the
softener to make up for sodium losses and to precipitate
excess calcium ion. The underflow is returned to the
reaction tank with the regenerated absorbent being pumped
to the absorber.
75
-------
6.2 PROCESS AND OPERATING CONSIDERATIONS
The effectiveness of a soluble alkali as an S0? scrubbing medium is
limited only by the gas/liquid chemical equilibrium and the rate of
transfer of SO^ from the gas to the scrubbing solution. Such scrubbing
systems are thus highly efficient in S02 removal. The main chemical
reactions involving absorption are:
2NaOH
The presence of oxygen in the gas stream will, as is common in all
absorption-based processes, convert the active sulfite species to inacti^e
sodium sulfate.
Na2S03 + 1/202 -»• Na2S04
2NaHS03 + 1/202 •+ Na^O^ + H20 + S02.
The concentration of the active alkali in a particular system determines
whether this process is described as the "dilute" or "concentrated"
system. In the dilute system, the active sodium is about 0.1 molar
while in the concentrated system it runs 0.5 molar or greater. There is
a marked difference between these two systems in their response to the
regeneration of the oxidized sodium sulfate product.
Ca(OH)2 + 2NaHS03
Ca(OH)9 + Na9SO. + 1/2H90 -»• 2NaOH + CaSO- 4- + H90
" " J fc J L.
Ca(OH)2 + Na2S04 + 1/2H20 -»• CaS04 • + 2H20+ + 2NaOH.
In the dilute system the Na^O^ is regenerated readily to NaOH under
appropriate operating conditions, and calcium sulfate is precipitated-
In the concentrated system, the above regeneration reaction does not
76
-------
TO STACK
S02 ABSORBER
GAS
CONDITIONING
SECTION |
I
I ,_J
NLET GAS FROM
DRY GAS CLEANING SECTION
fUME STORAGE
SODA ASH
STORAGE
t^VVVM j 1/V*VM—j
u
LIME
REACTION
TANK
CLARIFIER
WASH WATER
MIX
TANK
1
CLARIFIER
SLUDGE
TO DISPOSAL
TO VACUUM
PUMP
FILTRATE
RECEIVER
SURGE
TANK
SODIUM BASE DOUBLE ALKALI PROCESS-DILUTE SYSTEM
RGURE 6-1
-------
proceed readily and the soluble sodium sulfate tends to remain in the
system, necessitating a purge of the material with the associated problems
of potential water pollution or special treatment. However, in recent
pilot plant work with concentrated double alkali systems, Arthur D.
Little,Inc., has observed the simultaneous precipitation of sulfate and
sulfite; and it appears that if the system oxidation rate is below about
20-25 percent, sulfate can be removed without special purging and
treatment of the Na2SO,. Obviously further work is necessary before
these conclusions can be interpreted in terms of an actual commercially
operating process. At this point in the double-alkali process development'
in view of the overall regeneration chemistry, the dilute system appears
to be more suited to those applications where oxidation is expected to
be high, i.e., where the oxygen content of the effluent gases is
high.
In the dilute system, the precipitated sludge will contain an
appreciable proportion of calcium sulfate along with the calcium
Such sludges are less thixotropic, more easily filtered and can be more
completely dewatered than predominantly calcium sulfite sludges. These
are important characteristics if water conservation and a system water
balance is of concern. To minimize the loss of the active sodium ions»
washing of the sludge cake is mandatory; but at the same time, the
amount of washing must be limited if the system is to operate in the
closed loop mode and if the build-up of dissolved solids in the system
is to be controlled. Thus, the relatively coarse-grained calcium sulfat
crystals in the sludge provide favorable dewatering and structural
properties which in turn favor reduced washing cycles. It has been
reported that some double-alkali systems producing high calcium sulfate
2
cake can be filtered to over 65 percent solids.
Another important consideration in the double-alkali system is the
necessary softening step to control the level of dissolved calcium lot**-
remaining in the regenerated absorbent after the lime regeneration step*
If the regenerated absorbent is returned directly to the absorber,
i
is a potential of calcium sulfate or gypsum scaling within this equip
Sodium carbonate (Na-CO-) is usually used to provide both the sodium
makeup values and the carbonate for the softening process. The react!0
is:
+ Na2C03 •*• 2Na+ +
78
-------
If lime utilization in the regeneration step is poor, the carbonate
addition above may be inadequate to provide complete softening and the
use of C02 will be required to be added with the Na2CO. to remove the
excess calcium ions prior to returning the regenerated absorbent to the
scrubber. However, carbon dioxide also lowers the free hydroxide
Available for S02 absorption so the degree of softening by C02 should be
"eld to a minimum.
It should be noted that based on experience with the General Motors
full-scaie double alkali systems, carbonate scaling — i.e., the conversion
of soluble calcium ions to precipitated insoluble calcium carbonate in
absorbing system — can occur as a result of high pH (<9) scrubbing
or absorbing CO,, from the flue gas. Thus pH in the scrubbing system
8hould be controlled below this level.
As with all scrubbing systems, the liquid/gas ratio (L/G) in the
°uble-alkali systems directly affects S02 removal efficiency. For a
Siven removal efficiency, the L/G ratio must increase as the pressure
r°P in the absorber decreases. ' In the commercial scale system
nstalled at General Motors' Parma plant on coal-fired boilers, an L/G
tati0 of 20 gal/1000 CF is used for a column pressure drop of 7.5 inches
W r
Japanese installations on two oil-fired utility boilers use 12
8al/lOOO CF and 7.5 gal/1000 CF, respectively, at approximately the same
Pressure drops.5
Although in the U.S. regeneration has been affected through the use
lime, Japanese installations use limestone in the regeneration
ction of their concentrated double alkali systems. Preliminary laboratory
uting runs on dilute systems conducted by EPA and reported by Arthur
Little/EPA , together with rough economic analyses, suggest that the
Pital cost increases and increases in waste disposal costs for the
estone based regeneration system under U.S. conditions would exceed
thg
c°sts of the lime based systems. From the same studies, however,
e is confirmation that the use of limestone is a viable approach in
m
Degeneration of concentrated double-alkali absorbent stream.
PROCESS DEVELOPMENT
"he sodium base double alkali system has received considerable
ntion over the past several years both in the U.S. and in Japan.
°ugh there are no full-scale utility boiler applications in the
79
6 o
•'°
-------
U.S., there are two 150 MW operating systems in Japan on oil-fired
boilers. Both systems operate in the concentrated mode with gypsum
production. A total of 5 additional large scale utility concentrated
double-alkali systems with a total capacity of 1750 MW are presently
being engineered and/or constructed in Japan by Kawasaki/Kureka, with
start-up dates between mid-1975 and mid-1977.5 In the U.S. there are
two operating full-scale industrial boiler applications equivalent in
size to a 40 MW and a 20 MW power system, respectively, and an industrial
kiln control system. The 40 MW General Motors system is based on the
dilute process while the 20 MW Scholz station of Gulf Power Company
utilizes a concentrated system developed by Arthur D. Little/Combustion
Equipment Associates. The FMC concentrated system was put into service
in late 1971 to control the emissions from two reduction kilns. In
addition to these operating units, there are at least five full-scale
systems in the U.S. projected for completion by mid-1975. All these
units with one exception utilize the concentrated system.
It is obvious that in the application of the double-alkali process
to utility and industrial boiler installations, preference has been
directed towards the concentrated active alkali option. The usual
practice of operating industrial boilers with relatively high levels of
excess air would seem to suggest that special operating considerations
may be necessary with the concentrated system to optimize the regeneration
sequence. Operating experience and data analysis will be necessary
before the commercial performance of the concentrated system in a potential!
high oxidation rate application can be evaluated. Considerable development
work on both dilute and concentrated systems is still underway in the
U.S. by Environtech, General Motors, Zurn Industries and A.D. Little/
Combustion Equipment Associates; but there has been little pilot plant
work specifically directed towards the nonferrous smelter industry. The
Smelter Control Research Association (SCRA) has performed exploratory
investigations of both the double-alkali sodium and ammonia processes on
a 4000 SCFM pilot plant located at the McGill smelter of Kennecott
Copper Corporation. The dilute active salt mode was chosen but rapid
scale formation occurred in the scrubbing system since the softening
step was not apparently adopted. Some difficulty was also experienced
in regenerating the oxidized scrubbing liquor and it was concluded that
80
-------
the problems experienced made the process less attractive than the
ammonia double-alkali process, and attention has now been directed
accordingly to this process.
In spite of the impressive progress made in bringing the sodium
double -alkali process into commercial applications in the utility and
commercial boiler area, there are broad uncertainties regarding its
aPplication to dilute S02 streams containing relatively high oxygen
levels associated with reverberatory furnace operation in copper smel-
ters. Current programs sponsored by EPA, and continuing evaluation of
commercial scale units such as the General Motors Parma dilute mode in-
stallation and FMC's concentrated mode installations, should ultimately
resolve questions involving the effective regeneration of sodium sulfate
^ith good lime utilization and the potential use of limestone in the
regeneration cycle to improve overall economics.
5 '4 DESULFURIZATION EFFICIENCY
Performance of double alkali systems with respect to St^ removal is
ell established, with values of 90 to above 95 percent being common.
osorbent temperature, pH, concentration of the sulfite/bisulfite species
the total ionic strength all influence the equilibrium partial
e of S02 over solutions of sulfite/bisulfite and hence the S02
tem°val effectiveness. It is thus possible to design a double alkali
Astern to accomplish any desired level of removal efficiency.
In the Smelter Control Research Association's pilot plant work at
tK
6 Kennecott McGill Smelter on the double-alkali process, removal
*iciencies in excess of 90 percent were recorded with feed streams
C°ntaining 0.3 to 0.95 percent SO-. However, if the pH fell below 5.5,
etnoval efficiency fell sharply, with values of 55-75 percent being
tecorded.
*5 SULFITE OXIDATION
As has been noted in Section 6.2., the active alkali species in the
uole-alkali process — sodium sulfite and sodium bisulfite — are suscept-
ible *.
c to oxidation by the presence of oxygen in the gas stream, although
metallic particles or fume remaining after the gas cleaning
^Uence or impurities in the regeneration lime may also promote oxidation.
81
-------
The rate of oxidation is thus related to the composition of the
scrubbing liquor (high concentrations of dissolved salts may absorb
oxygen more slowly), the oxygen content of the gas, the presence of
impurities and the equipment design itself which may limit oxygen uptake.
The sodium sulfate produced by the oxidation reactions is inactive
in the SCL absorption process and the overall result of oxidation is to
remove active sodium from the scrubbing circuit. This soluble sulfate
must be removed from the system but there are certain problems which
must be recognized:
1) A simple purge of the soluble Na-SO, to disposal poses
environmental problems
2) A large loss of sulfate means active sodium must be replaced
in the scrubbing circuit and this imposes a substantial
economic penalty.
From a practical point of view, it is necessary to convert this
sulfate back into active sodium by removing the sulfate ion. Several
techniques are available.
In the dilute double-alkali system, treatment with lime as noted in
Section 6.2, removes the sulfate ion as gypsum (CaSO, '2H..O) with the
production of NaOH. In Japan, where a market for gypsum exists, the
sulfate ions are precipitated from the concentrated double-alkali solutions
by acidulating with sulfuric acid. The applicable equation is:
Na0SO, + 2CaSO_-l/2H00 + H.SO, + 3H00 -> 2NaHS00 + 2CaSO, -2H00
^.H J/^4Z 3 42
This approach is not economically attractive if the gypsum must be
discarded and where the oxidation rate is high. In experimental work
conducted by Arthur D. Little under the EPA-ADL Dual Alkali Program, it
was observed that a simultaneous precipitation of sulfate and sulfite
with lime and limestone treatment of concentrated double alkali liquors
took place if the system oxidation rate was below about 20-25 percent of
the SO^ removed. As noted earlier, this phenomenon suggests the possibility
of a broader potential for the concentrated double alkali system.
82
-------
It is also possible to limit the degree of oxidation through process
and equipment design. Minimum residence times and the use of high ionic
strengths in the scrubbing liquor provide two such approaches. The use
°f anti-oxidants may also be useful although such compounds are usually
e*pensive and add appreciably to the operating costs.
6-6 FEED GAS TREATMENT
Since the double-alkali process produces a throwaway product, it is
n°t necessary to provide the same degree of solids cleaning as that
Required for a closed loop regenerative system, and gas pretreatment
requirements are satisified by cooling the feed gas stream to around
30°F. A system similar to that discussed in Appendix A but without the
Provision of mist electrostatic precipitators should provide acceptable
c°oling and saturation of the gas with moisture to minimize evaporation
and scaling in the absorber itself.
6'7 PARTICULATE REMOVAL CAPABILITY
The removal of particulates in the scrubbing operation of the
°uble-alkali systems is variously reported. This is to be expected
ince removal efficiency in scrubbing is affected by a large number of
ariables such as liquid to gas ratios, particle size distribution,
ature of the particles, energy input to the scrubber, and type of
Crubber. Where no particulate removal is needed, as in the case of a
*e~cleaned gas stream, simple baffle-type scrubbers can be used to
em°ve SO-. Where simultaneous particulate and,S02 removal is necessary,
8h energy venturi scrubbers can effectively do both jobs.
Although a scrubber can be designed to achieve any reasonable
Qa _
8tee of particulate removal, depending basically on the amount of
Tjirj _
eumatlc and hydraulic energy input at the contacting surface, SO.,
Jlh
s°rption columns are not high pressure drop devices and they are not
rti-cuiarly effective particulate removers. As has been noted in
ction 2, gas cleaning or particulate removal has been considered as an
ePendent process step which can be tailored to suit the requirements
of t,
ne SO- absorber section and the condition of the incoming gas. In
ters, the use of cyclones, balloon flues and electrostatic pre-
dators is universal on reyerberatory gas streams, and the mass
•^S of the residual particulates is low with the size distribution
-------
6.8 PROCESS ENERGY REQUIREMENTS
The double-alkali process is not an energy-intensive process.
Electrical power requirement to support the absorption process, liquor
circulation, and the solids separation equipment is a function of the
total gas flow and the S02 handling rate in the regeneration section;
but total power costs amount to only about 6-8 percent of the total
direct annual operating costs for a range of gas flows from 70,000 to
300,000 SCFM containing 1 percent S02.
6.9 RETROFITTING REQUIREMENTS
There is some potential of separating the gas conditioning and S02
absorption section of the process from the regeneration section in this
process so retrofitting will predominantly affect the location of the
gas conditioning equipment and the S02 absorber in relation to the
existing ductwork and available space. Space requirements for this
equipment are similar to those required for the other absorbent-based
systems. With the exception of the clarifiers, the tankage and other
equipment associated with the regeneration section can be installed in
a number of ways. The equipment at General Motors Parma plant occupies
three floor levels, all under cover, Another vendor of the double-
g
alkali process, FMC, estimates that the recovery plant for a 50 MW
power plant (approximately 100,000 SCFM) would occupy two floors, each
40' x 60'.
Generalized space requirements can be estimated as follows:
Gas conditioning and S02 absorption: 8-12,000 sq. ft.
(including neutralization and cooling
tower)
Regeneration section and clarifiers 20-25,000 sq. ft.
-------
6.10 COSTS
8 9 10
A number of cost estimates ' ' for the double alkali process have
been published, and although they cover both the dilute and the concen-
trated absorbent system modes, they have provided source data in deve-
loping both capital and operating cost relationships for the S0_ handling
section of the process. A capital cost curve based on the hourly rate
°f SO- handled by the regeneration section is provided by Figure 6-4.
Capital costs for the S02 absorption section including absorbent re-
circulation have been taken from the appropriate generalized curve for
clear solution absorption in Appendix B. Figure 6-3 represents the
combination of these two cost sections under defined conditions of gas
flow and S02 content.
The direct annual operating costs were developed from reported
usage values and updated unit costs presented in Table 6-1. The corres-
ponding direct annual operating costs for the range of S02 contents are
Provided in Figure 6-3. It should be noted that these costs do not
Delude capitalization costs.
Tables 6-2 and 6-3,respectively, provide capital and operating cost
etails for the double-alkali process itself and for the overall control
Astern which includes the gas conditioning section. These tables provide
ir
comparison with similar tables for the other S0» control processes
Utl
-------
5) In the dilute system, the sulfate oxidation product is
regenerated during the lime treatment step with precipitation
of calcium sulfate together with the calcium sulfite.
Disadvantages are:
1) Supplemental "softening" of the absorbent stream is
required to control residual calcium sulfate and potential
scaling in the absorber system.
2) The double alkali system shares with the lime/limestone
systems the disadvantage of producing large volumes of waste
calcium sulfite/sulfate (3-4 lbs/lbS02) which must be disposed
of in an environmentally acceptable manner.
3) Soluble sodium salts will end up in the waste sludges and
although the level should be low, they may pose water
pollution problems.
6.12 REFERENCES
1. LaMantia, C., et al., "EPA-ADL Dual Alkali Program Interim Results."
Presented at EPA Symposium on Flue Gas Desulfurization, Atlanta,
Ga. Nov. 4-7, 1974.
2. Cornell, C.F., "Liquid-Solids Separation in Air Pollutant Removal
Systems." Presented at the ASCE Annual and National Environmental
Engineering Convention, Kansas City, Missouri, Oct. 1974.
3. Cornell, C.F., Dahlstrom, D.A., "Performance of a 2500 ACFM Double
Alkali Plant for S02 Removal." 66th Annual Mtg., A.I.Ch.E., 1973.
4. Asahara, Ken, "Double Alkali Systems for Control of Sulfur Dioxide
Pollution." Chemical Economy and Engineering Review, December
1972.
5. Kaplan, Norman, "An Overview of Double Alkali Processes for Flue
Gas Desulfurization," Presented at EPA Symposium on FGD, Atlanta,
Ga., Nov. 1974.
6. Phillips, R.J. "Operating Experiences with a Commercial Dual-
Alkali S0? Removal System." Presented at 67th Annual Mtg. of APCA,
Denver, Col., June 1974.
7. Campbell, I. E., "Exploration Investigations of the Ammonia and
Sodium Double Alkali Processes for S0~ Control." Presented at
AIChE National Mtg., Salt Lake City, Utah, August 1974.
8. Brady, J.D.. "FMC Corporation's Sulfite Absorption-Lime Regeneration
Process." Presented at Missouri Public Service Commission Conference,
Sept. 1974.
86
-------
Smelter Control Research Association, Inc. "Report to U.S. Bureau
of Mines on Engineering Evaluation of Possible High Efficiency
Soluble Scrubbing Systems for the Removal of SCL from Copper Smelter
Reverberatory Furnace and Like Flue Gases," March 1974.
G.G. McGlamery and R.L. Torstrick, Tennessee Valley Authority,
"Cost Comparisons of Flue Gas Desulfurization Systems." Presented
at FGD Symposium, Atlanta, Ga., Nov. 1974.
87
-------
Table 6.1. SODIUM SCRUBBING - DOUBLE ALKALI (THROWAWAY)
UNIT USAGE AND COST DATA
A. Chemicals & Utilities
Basis
Unit Cost
Lime
co2
Power a) Absorption
b) Regeneration
Water
1.25 lb/lbS02
0.038 lb/lbS02
0.0116 Ib/lbSO,
i
4KW/1000 SCFM
0.075 KWh/lbSO,
0.2 gal/lbS02
$22/ton
$52/ton
$50/ton
$0.015/KWH
$0.015/KWH
$0.30/M gal
B. Operating Labor &
Maintenance
Labor a) Absorption
b) Regeneration
Maintenance
Taxes & Insurance
Sludge Disposal
% man/shift for
<75,000 SCFM
3/4 man/shift for
>75,000 SCFM
1 man/shift
<10,000 lbS02/hr
lh man/shift
>10,000 lbS02/hr
4% TCI/yr
2 1/2% TCI/yr
3.53 Ib Sludge/lbS02
$8/hr
$3/ton
C. Fixed Charges 13.15% TCI/yr
Based on Capital Recovery Factor using 10% interest over 15 year life
-------
DOUBLE ALKALI PROCESS -SODIUM BASE
TOTAL CAPITAL INVESTMENT COSTS
S°2 REMOVAL EFFICIENCY 95%
$10-0 Million
NOTE:
GAS COOLING & CONDITIONING
NOT INCLUDED.
4.0% SO,
0% SO,
0.5% SO-
1974 Costs
100,000 SCFM
GAS FLOW RATE-SCFM
89
FIGURE 6-2
-------
DOUBLE ALKALI PROCESS - SODIUM BASE
TOTAL DIRECT ANNUAL OPERATING COSTS
S02 REMOVAL EFFICIENCY 95%
4.0% SO,
$10 Million
NOTE: GAS COOLING & CONDITIONING
NOT INCLUDED.
•^^••••••••••••••••••••••^^•••^•••••i^^PM^"*^^*^^^""*"**
1.0 Million
2.0% SO,
1.5% SO,
1.0% SO
100,000 SCFM
GAS FLOW RATE-SCFM
90 ,
-------
SODIUM DOUBLE ALKALI SCRUBBING
TOTAL CAPITAL INVESTMENT COSTS
REGENERATION & PRECIPITATE HANDLING
$10.0 Million
Costs: Mid 1974
10,000 LBS/HR
SULFUR DIOXIDE RATE
OF REGENERATION SYSTEM (Ibs/hr)
91
FIGURE 6-4
-------
Table 6.2. SODIUM SCRUBBING - DOUBLE ALKALI DILUTE SYSTEM (THROWAWAY)
CAPITAL AND TOTAL ANNUAL COST
vo
NJ
TOTAL CAPITAL INVESTMENT
Annual Cost
A. Direct Operating
1. Power
2. Water
3. CaO, Na2C03, C02
4 . Labor
-*• Maintenance
6. Sludge Disposal
7. Taxes & Insurance
TOTAL ANNUAL DIRECT
OPERATING COST
B. Annual Capitalized
Cost
C. Net Total
Annualized Cost*
*BaseA. OTL CorvoTate tax. \Lat
$
$/SCFM
$/ Annual ton
S02 Removed
$/yr/ton
S0» Removed
$/yr/,ton
SO 2 Removed
e o£ «v8%.
(
70,000
4,200,000
60
152
96,000
3,300
846,100
100,000
168,600
290,000
105,000
1,609,600
59
554,400
1,257,000
46
3as Flow Rate - !
100,000
5,250,000
53
134
137,200
4,700
1,212,300
151,200
210,000
411,500
131,300
2,258,200
58
693,000
1,699,000
43
3CFM @1% S02
200,000
8,100,000
41
103
274,800
9,400
2,424,600
151,200
324,000
831,900
202,500
4,218,400
54
1,069,200
3,004,000
38
300,000
10,750,000
36
91
411,600
14,100
3,637,200
151,200
430,000
1,235,000
268,800
6,147,900
52
1,419,000
4,272,000
36
-------
Table 6.3. SODIUM SCRUBBING ~ DOUBLE ALKALI DILUTE SYSTEM (THROWAWAY)
TOTAL SYSTEM CAPITAL AND ANNUAL OPEBATING COSTS
A. TOTAL CAPITAL INVESTMENT
1. Gas Conditioning
2. Absorption and
Neutralization
TOTAL
B. TOTAL ANNUAL
OPERATING COSTS
1. Gas Conditioning
2. Absorption and
Neutralization
TOTAL
C. NET TOTAL
ANNUALIZED COST*
1. Gas Conditioning
2. Absorption and
Neutralizat ion
$ /Annual ton
S02 Removed
$/yr/ton
S02 Removed
$/yr/ton
S02 Removed
(
70,000
1,210,000
4,200,000
5,410,000
197
211,100
1,609,600
1,820.700
67
230,200
1,257,000
1,487,200
54
Jas Flow Rate -
100,000
1,500,000
5,250,000
6,750,000
172
287,000
2,258,200
2,545,200
65
298,500
1,699,000
1,997,500
51
SCFM @1% SO-
200,000
2,250,000
8,100,000
10,350,000
132
465,200
4,218,400 .
4,683,600
60
465,800
3,004,000
3,469,800
44
300,000
2,900,000
10,750,000
13,650,000
112
635,700
6,147,900
6,783,600
58
619,200
4,272,000
4,891,200
42
*Based on Corporate Tax Rate of 48%.
-------
94
-------
7.0 MAGNESIUM OXIDE SCRUBBING
7-l PROCESS DESCRIPTION
There are a number of different magnesia-based scrubbing systems
which provide effective S02 removal, but U.S. development work, as well
as Russian and Japanese, has concentrated on the use of magnesium sulfite-
^gnesium oxide slurries having a basic pH. This process was the one
chosen for the demonstration installation at Boston Edison's Mystic
"tation. The process description provided below and represented in
igure 7.1 is based on this approach but includes the process alter-
natives proposed by TVA in their conceptual design and cost study on
scrubbing for the slurry handling and S02 recovery sections.
are four primary operations:
1) Absorption
2) Slurry handling
3) Drying - Calcination
4) MgO system
Absorption — The S0~-containing gas streams, after suitable cooling
removal of particulate matter, enters a venturi or absorber tower
e
e it is scrubbed with the magnesia slurry.
Blurry Handling — A bleed stream is taken from the slurry re-
_ . " • ..... • ••- .— •Jfc
rculation loop directly to a centrifuge to provide a wet cake of
8 °3*6H 0 and MgSO, or, alternatively, to reduce overall energy require-
^ntp
to gravity thickeners or wet screens for thickening the solids,
nve*sion of the MgSO»'6H70 to MgSO~-3H20 by heating of the slurry, and
fchen t-n
to centrifuging.
." • Sgyjjig - Calcination ~ The wet cake of MgS03'3H20 with some MgSO^
rted in a fluid bed dryer at approximately 400°F to a moisture
ent of less than approximately 4 percent. Off gases are cleaned by
UB
U8e of cyclones and a high temperature bag filter or ESP. The dried
•Lin
with added coke to reduce the MgSO, are conveyed to a fluid bed
^Cin
j ner operating at 1400 to 1600°F. Provision of a waste heat boiler
e °ff gas stream to cool to 600-700°F followed by air dilution to
t
0 400°F allows use of a bag filter for dust control. The clean
95
-------
and diluted S02 gas stream containing approximately 8 percent SCL is
directed to a sulfuric acid plant or an elemental sulfur plant.
MgO System — The regenerated MgO is conveyed to storage and with
the required makeup is reslurried for circulation to the absorption
section.
7.2 PROCESS AND OPERATING CONSIDERATIONS
Slurries of magnesia are good absorbants for SO- and the process
provides easy separation of the sulfite salts from the scrubber liquor,
an ability to regenerate and recycle the absorbent, and an avoidance of
a solid disposal problem. The main reactions which take place when the
S02~containing gas is contacted with an aqueous recycled slurry of
magnesium oxide (MgO), magnesium sulfite (MgSOO and magnesium sulfate
(MgS04) are: ;
Main Reaction: MgO + S02 + 3H20 -» MgS03«6H20
Side Reactions: MgS03 + S02 + H-O* Mg(HSO )
O <£
Mg(HS02)2 + MgO + 2MgSO + H,0
J £
MgO + S03 + 2H20 -> MgS04- 2H 0
MgS03 + 1/2 02 -*-MgS04- 2H20
The important relationships in the MgO slurry process which directly
affect the degree of S02 removal and thus the operating design of the
process are:
1) Liquid to gas ratio (L/G) - defined as the
ahsnr'hpnl- npf 1 flflfi ATFM on -.
2S02 reLcbt l^J^*~ **?"" « -creased
strong influence on the S00 vapor J?i« " coraP^ition exercises a
scrubbing efficiency is direcSy relaJeS to^h" J?I/0lUtl0S ^
the partial pressure of SO in the «S!v ^difference between
pressure over the scrubbing solution SQ8"8 "* *" ^ Vap°r
with increased PH. option. S02 vapor pressure decreases
96
-------
TO STACK
GAS
'CONDITIONING
SECTION
. J
INLET SMELTER GAS
ABSORBER
(VEWTUR/J
WETYV
SCREENN\
V~
PURGE
COLLEC
CONVERSION
irTANK
COKE
STORAGE
CONVEYOR- - ELEVATOR
FLUID BED
AIR DRYER
OIL
FLUID
BED
CAUCINER
MgO
MAKE UP
MgO
STORAGE
"^RECYCLE
MgO
STORAGE
SLURRY
TANK
WASTE
BAG
FILTER
TO SULFURIC ACID PLANT
MAGNESIUM OXIDE PROCESS
FIGURE 7.1
-------
Temperature exercises little adverse effect on the mass transfer
rates with a pH above 6 because of this very low S02 vapor pressure over
the scrubbing solution at this pH level. The effect of MgSO^ on S02
absorption is also minor at higher pH values.
Both venturi and mobile bed absorbers have been evaluated for
magnesia scrubbing and both types are capable of attaining S02 absorption
efficiencies of 90 percent or greater, but the venturi absorber requires
a higher operating L/G to achieve the same absorption efficiency. In
spite of this factor, the venturi absorber appears to provide overall
operating advantages and it has been selected in one configuration or
another as the SOp absorber for the commercial U.S. installations.
Typical operating parameters of a venturi absorber scrubbing utility
plant flue gas are:
L/G: 20 gal/MACFM
Gas Velocity: 75 ft/sec
SP: 4-6" WG.
pH: 7-7.5.
On the basis of Chertkov's work, SO^ concentration in the gas phase
would not be expected to adversely affect the absorption rate until
concentration levels are above 3.5 percent, but as S0~ concentrations in
the stack gas increase and magnesia requirements increase correspondingly*
the L/G gas ratio will also increase to maintain the usual 10 percent
weight of solids in the slurry.
Because the magnesia scrubbing system operates in a closed loop,
consideration must be given to the buildup of impurities in the system
and the requirement for purging. The quality of the make-up water, the
impurities of the magnesium oxide itself, together with the level of
soluble MgSO,, determine the purge required; and if losses of active
salts are to be minimized, these elements must be controlled. In the
case of magnesium oxide, this will limit the sources of MgO suitable for
magnesia scrubbing. Raw uncalcined magnesite (45 percent MgO) which
occurs in Nevada and agricultural-grade calcined magnesite (87 to 92
percent MgO) would increase the purge requirements significantly. The
requirement for 98 percent MgO calcined magnesite may involve not only
initial cost premiums but appreciable transportation costs as well,
depending on source and use location. Unusual local water supply
may necessitate additional pretreatment before introduction into the
magnesia scrubbing loop.
98
-------
The process as now commercially defined imposes no unusual problems
in terms of corrosion or special processing equipment, although certain
processing areas appear to offer potential for further processing
development. As is normal in slurry systems, the use of rubber lined or
specially coated equipment and piping in the scrubbing and recirculation
system has provided acceptable service in the U.S. commercial installations
apart from some localized problem areas. In the Boston Edison installation,
the wet solids handling section, centrifuge operation, and wet cake
handling provided difficulties during operation which may suggest an
area for further development. The use of fluid bed equipment for drying
and calcination in place of the rotary equipment used is an attractive
alternative which offers the potential of reduced operating costs and
improved process control.
The S0? removed in the calcination step may be directed after
Suitable gas cleaning to either a conventional sulfuric acid plant or an
elemental sulfur plant. The nature of the recovery process also provides
the further option of separating the S02 regeneration and utilization
steps from the gas scrubbing and slurry processing section. The U.S.
commercial demonstration installations of the magnesia scrubbing process
have taken this separate S02 regeneration and use approach using a
8ulfuric acid plant. The concept of a centralized regeneration and
Processing facility servicing a number of independent scrubbing facilities,
however, may offer significant economies in a situation where different
scrubbing facilities are within the same geographical area and a market
e*ists for the sulfuric acid produced.
'3 PROCESS DEVELOPMENT STATUS
The magnesium oxide slurry scrubbing system has been under commercial
8cale evaluation in the United States since 1972.
Construction of the first commercial system, the Chemico/Basic MgO
sulfur recovery processj including the scrubbing and recovery system at
B°ston Edison's Mystic Station on a 150 MW oil-fired boiler, and the
Generation facility at the Essex Chemical Company, Rumford, R.I.,
8ul£uric acid plant, was completed in April 1972. A second Chemico/Basic
Astern Was piaced ln operation in September 1973 at Potomac Electric
^
°Wer Company's Dickerson Station on half the flue gas from a coal-fired
rated at 190 MW. A third system using magnesium-sulfite slurry
99
-------
scrubbing is currently being installed on an equivalent 120 MW coal
burning boiler at Philadelphia Electric Company's Eddystone Station.
The Boston Edison system during its 27-month period of operation
(it was shut down in June 1974 at completion of the contract) logged
4000 hours of running time and demonstrated that it could meet the
process guarantees of 90 percent removal of the inlet S02, that the
magnesia could be regenerated and recycled, and that 98 percent sulfuric
acid of good quality could be recovered from the S09 removed from the
4 z
flue gas. Equipment problems and malfunctions were frequent during the
test period, especially in the dryer and calciner systems, but they were
corrected before completion of the program. These problems also contributed
to the higher MgO losses experienced during the program (10 percent) and
prevented demonstration that the regenerated MgO could be recycled
continuously without loss of reactivity and determination of the effect
of buildup of particulate matter, vanadium, etc., in the system. The
program did provide additional valuable information for future plant
design in such areas as required quality control and analytical methods,
equipment selection, and materials of construction.
The buildup of impurities, both soluble and insoluble, introduced
into the closed loop system by makeup water, makeup MgO, oxidized
MgSO, , and particulates captured in the S00 absorber will require some
^ £•
attention in an operating facility, but the limited operating experience
gained to date does not allow a clear definition of the magnitude of the
problem and the appropriate approach required. In dry climates, with
high solar evaporation rates, bleeding off a small stream to a dead-end
pond may be acceptable though losses of magnesium via MgSO. must be
considered.
The two process sections which appear to offer potential for improve*
reliability and/or reduced energy requirements are:
1) The slurry or solution process section
2) The drying and calcining sections.
In the MgO slurry scrubbing system, the effluent from the S0
absorber contains predominantly MgSO~*6H90. Private in-house and
1
communications reported by TVA indicate that MgSO-'6H90 rapidly convert3
•J £
to MgS03'3H20 at the reasonably low temperature of 80°C and that the
conversion rate is significantly increased by increasing the slurry
concentration from 10 to 60 percent MgSO-. Although dewatering of the
100
-------
MgSO_*3H20 may be somewhat more difficult because of the smaller crystal
size, less energy is required to dehydrate the trihydrate than the
hexahydrate, and this approach may provide a viable route to reduced
operating costs. The TVA report referenced above provides a number of
separation alternatives with associated capital and operating costs
related to a power plant facility, but the relationships have general
application.
In the drying and calcining sections, the use of fluid bed equip-
ment offers advantages in lower investment and operating costs, lower
heat losses, and better temperature control and in the case of the
calciner, precise control of oxygen which could reduce the requirement
for coke to reduce the MgSO,. It should be noted that the regeneration
facilities for the Philadelphia Electric Company's Eddystone magnesia
scrubbing system will include a fluidized bed reactor. On the basis of
Presently operating and planned installations, the magnesia scrubbing
Process appears to have qualified as a demonstrated commercially viable
desulfurization process. In addition to the three U.S. installations
n°ted above, Philadelphia Electric Company has signed an agreement with
E**A committing it to the installation of additional magnesia scrubbing
systems by 1978 at their Eddystone 1 and 2 stations and their Cromby-1
Station. A number of installations are reported in operation in Japan.
TT
w° units are installed in copper smelters and are tied to sulfuric acid
Plants; one treats 44,000 SCFM of tail gas from a sulfuric acid plant
and the second treats 49,000 SCFM of converter gas containing 2 percent
°2" Another installation to start up in the near future will treat
5,000 SCFM gas from a Glaus furnace and an industrial boiler to produce
eleroental sulfur via the Glaus unit.
•* PROCESS DESULFURIZATION EFFICIENCY
Commercial demonstration runs on both coal- and oil-fired utility
oilers have demonstrated that the magnesium oxide scrubbing process has
n s°2 removal capability on these streams up to at least 90 percent.
°nsiderable experimental work over several decades has established with
1116 assurance the chemistry, kinetics and mass transfer relationships
Of j.,
cne different magnesia scrubbing systems; and with appropriate adjustment
101
-------
of the L/G ratio and control of pH, it appears that this level of S02
removal efficiency can be maintained using magnesium oxide slurries on
SCL-containing gas streams containing up to 3 to 4 percent S0?. However,
magnesium sulfite-bisulfite scrubbing, because of the higher equilibria
vapor pressure of the S0? over a slurry of MgSCL in a slightly acid
solution of bisulfite, may not be quite as effective as the alkaline
scrubbing system in SO,., removal efficiency.
7.5 SULFITE OXIDATION
Both the MgO and MgSO~ species in the scrubber slurry are susceptible
2
to oxidation to the soluble MgSO,. Downs and Kubasco indicate that
most of the magnesium sulfate formed in mobile bed or venturi scrubbers
results from sulfite oxidation by oxygen absorbed from the input gas. A
number of researchers have also noted the catalytic effect of heavy
metal ions on increased sulfite oxidation rates. Chertkov found,
however, that the concentration of the oxidation product itself, MgSO,,
has a significant influence on reducing oxygen mass transfer coefficients
and presumably sulfite oxidation rates. Higher pH also reduces oxidation-
Other work by Chertkov suggests that the oxidation rate above 13 to 15
percent MgSO, in the slurry may be small or even zero. If the steady-
state oxidation rate is not zero, provision must be made for MgSO,
removal. Data from the Boston Edison demonstration indicates that some
MgSO, (approximately 4 percent) is occluded in the MgSO_»6H00 crystals
ft
and removed with the centrifuge cake. If the MgSO-*6H_0 to MgSO»*3H90
J 2. j *>
conversion route is adopted, the occluded MgSO, will be solubilized and
returned with the mother liquor to the scrubber system. In this case
separate purge facilities for MgSO, may have to be provided. Although
there is little information currently available to predict the effect of
long term buildup of MgSO, in continuously operated, closed cycle S02
scrubbers, operating experiences reported by Chemico on the Boston
4
Edison installation suggest that changes in the S02/02 ratio with
reduced S02 causes higher oxidation rates with the increased MgSO,
concentration making centrifugal separation more difficult. Where the
gas stream is likely to contain high oxygen levels such as copper rever-
beratory flue gases, the use of organic inhibitors may be warranted to
reduce their oxidation rate.
102
-------
Some oxidation of MgSO- to MgSO, may also take place during the
drying cycle but complete regeneration to MgO is readily achieved in the
calciner with the addition of coke.
7.6 FEED GAS PRETREATMENT
The magnesia scrubbing process in common with most aqueous scrubbing
systems requires the feed gas to be at least saturated with water vapor
to minimize evaporation and localized high salt concentrations in the
absorber. With initial gas temperatures of between 300 to 600°F, water
Quenching in the typical system discussed in Appendix A should provide
acceptable conditioning. The closed system mode of operation of the
^agnesia process emphasizes the advantages of minimizing the intro-
duction of both particulates, particularly oxidation catalyzing metals,
and sulfuric acid mist into the S02 absorber and the gas pretreatment
Section should include an electrostatic mist eliminator or similar
device.
7'7 PARTICULATE REMOVAL CAPABILITY
Particulate removal efficiency is related to the nature of the
Particles themselves, their mass loading, particle size-and distribution,
arid the pressure drop of scrubbing equipment. However, in discussing
"e Particulate removal capability of the magnesia scrubbing system,
attention will be focused only on the S02 absorber itself and the response
wl»ich might be expected when treating a gas already cooled and conditioned
by 3
a system such as that described in Appendix A.
The venturi scrubber commonly used as the S02 absorber in magnesia
systems is not a high pressure drop device at 4 to 6" W.G.
W1 #-t_
h electrostatic precipitator cleaning of the gas stream prior to
nditioning, and with the application of a mist precipitator after
nditioning, the particulate loading into the absorber will be low,
^i t*i*
the size distribution itself concentrated under 2 to 3 y. Further
PatM
iculate removal in the scrubber will be limited.
, A test conducted by York Research Corporation on the venturi S02
fe °tber installed at the Boston Edison Mystic Station on an oil-fired
er indicated an average mass removal efficiency of around 57 percent.
103
-------
Tests conducted to determine efficiency of removal according to particle
size indicated that approximately 80 percent of the particulate material
in the inlet gas was sized greater than 7 y • About 10 to 11 weight
percent of the incoming particulates were sized at and below 1 y with
around 50 weight percent of this material being removed in the scrubber.
With input gases which have already received particulate cleaning, and
having a remaining particulate distribution predominantly under the 2 to
3 y level, the particulate removal efficiency in the scrubber would be
expected to be well below 50 percent. The nature of the particles
themselves, poor wettability, etc., may also increase the difficulty of
capture in the absorber.
7.8 PROCESS ENERGY REQUIREMENTS
The overall energy requirements for the magnesia scrubbing system
are determined largely by the concentration of the S02 in the treated
gas streams. Power requirements for the gas conditioning system, and
moving the gas stream through the SO- absorber, demister and ductwork
will be essentially the same as for other desulfurization processes,
although column design and related pressure drop will introduce some
differences.
In the slurry handling and S02 recovery sections, the S02 rate will
directly affect the power requirements of the centrifuges, conveyors,
blowers to the driers and calciners, and slurry preparation area; but a
requirement of 0.1 to 0.2 Kw/lb S02 recovered appears likely.
The most significant energy demand is imposed by the drying and
calcining steps of the MgO regeneration process. Direct firing of
either fluidized or rotary units requires either gas or oil fuel. The
Boston Eddison-Essex Chemical demonstration unit used No.6 fuel oil to
fire both the rotary drier and the rotary calciner. An approximate
estimate of the heat requirements for such a system indicates a value of
7000 to 8000 Btu/lb S02 or 0.05 gallons of fuel oil per Ib S02 recovered.
At today's oil prices, the cost of S02 regeneration in the magnesia
scrubbing process becomes the most significant element in the overall
operating cost picture. Even if fluidized beds are used and advantage
is taken of the lower drying requirements of the MgSO~* 3H20 crystal,
oil requirements still appear to be on the order of 0.04 gal/lbS02
104
-------
recovered. Fluidized bed drying and calcining direct annual costs for a
gas flow of 100,000 SCFM containing 1 percent S02 are approximately 60
Percent of the total annualized operating costs including capitalization
costs. As the concentration of S02 in the input gas goes up, the S02
recovery energy costs take an increasingly larger share of the total
annual operating costs.
7.9 RETROFITTING REQUIREMENTS
The magnesia scrubbing process as a whole is readily separated into
three separate areas, and direct retrofitting considerations apply only
to the gas conditioning/SO™ absorption section. The slurry processing
and SO- recovery area, and the S02 conversion section—sulfuric acid
Plant or elemental sulfur plant—can be located apart from the gas
handling section and even apart from each other, and this situation
allows considerable flexibility in adjusting to existing plant con-
figurations.
Space requirements for the separate areas of the process can vary
widely depending on equipment options such as rotary or fluid bed
drying equipment. Throughput, in terms of both gas volume handled and
Weight of equivalent S02 recovered, will have some effect on space require-
ments; but within the ranges covered in this study, the effect is not
e*pected to be large.
Space requirements are estimated as follows:
Gas Conditioning and S02 Absorption: 8 to 12,000 sq, ft.
Neutralization and Cooling Tower: 2 to 3,000 sq. ft.
Slurry Processing and S02 Recovery:
(fluidized handling) 25 to 35,000 sq. ft.
(rotary handling) 50 to 60,000 sq. ft.
H2SO^ Plant including 30 Day Storage: 35 to 45,000 sq. ft.
Sulfur Plant 15 to 20,000 sq. ft.
105
-------
7.10 PROCESS COSTS
The Tennessee Valley Authority's conceptual design and cost study
provided capital and operating cost estimates for this process as applied
to various sized utility plants. Figure 7-4 provides a capital cost
curve for the S02 regeneration section which reflects the TVA cost
estimates for this section with appropriate escalation and indirect
changes. The capital costs for the S0« absorption section have been
taken from the general cost curves developed in Appendix B. These two
sets of data have been combined to develop the parametric capital cost
curves relating SO,, concentration and gas flow rate provided in Figure
7-2.
Operating costs which also cover both the absorption section and
the regeneration section have been developed from data reported in the
TVA report above and in a summary of operating results from the Boston
Edison installation. Specific values are provided in Table 7.1.
Total annual direct operating cost curves for a range of S0« concen-
trations are provided in Figure 7-3.
Tables 7.2 and 7.3, respectively, provide capital and operating
cost details for the magnesium oxide process itself and for the overall
control system, which includes the gas conditioning section and an
auxiliary sulfuric acid plant. These tables provide direct comparison
with similar tables for the other S02 control processes under evaluation.
7.11 ADVANTAGES AND DISADVANTAGES
Advantages of this process are:
1) Process has been commercially demonstrated in both the
the United States and Japan and additional units are
presently in construction and/or planning stages in the
United States. Long-term reliability can be expected
to continue to improve quite rapidly.
2) Utilization of the regenerated S02 can be obtained as high
grade sulfuric acid via a conventional H,,SO, plant or as
elemental sulfur via an S02 reduction step.
3) There is an appreciable degree of independence between the
regeneration step and the SO,, removal step. Outages of the
regeneration facilities and the end use acid or sulfur plant
can be tolerated without interruption of the SO- section.
106
-------
This flexibility also allows the regeneration step to be
accomplished at a different location from the S09 control
point and offers the potential of having one regeneration
and SO™ use facility servicing a number of S0» control
facilities. l
4) The use of the venturi scrubber for S02 removal provides
considerable turn-down capability and the ability to handle
fluctuating input gas flows without incurring problems either
within the scrubber itself or in the slurry handling section.
5) The process produces only minor quantities of waste material
associated with the purge requirements of the systems.
The most important disadvantage of the process is:
1) The energy requirements for S02 regeneration are high and
can account for 60 percent or more of the total annual
costs. They are directly related to the SCL content of the
gas to be treated and the SO- recovered. Although the SO
removal capability of the process is excellent over a wide
range of $©2 contents, the related energy requirements suggest
that its application potential is best utilized on low S09
content (< 2 percent SO. gas streams).
7-12 REFERENCES
I* McClamery, G.G., Torstrick, R.L., et al. "Conceptual Design and
Cost Study - Sulfur Oxide Removal from Power Plant Stack Gas,"
Tennessee Valley Authority, EPA-R2-73-244, May 1973.
2- Downs, W., and Kubasco, A.J., "Magnesia Base Wet Scrubbing of Pul-
verized Coal Generated Flue Gas - Pilot Demonstration," Babcox and
Wilcox Company for EPA, Report 5153, September, 1970.
•*• Chertkov, B.A. "The Influence of S02 Concentration in a Gas on its
Rate of Absorption by Different Solvents," Khim. Prom. 7. 586-91,
1959.
*• Koehler, G.R., and Dober, E.J. "New England SO^ Control Project-
Final Results," Chemical Construction Company, Paper Presented
at EPA Flue Gas Desulfurization Symposium, Atlanta, Georgia, November,
1974.
Jumpei, Ando. "Utilizing and Disposing of Sulfur Products from FGD
Processes in Japan." Paper Presented at EPA Flue Gas Desulfurization
Symposium, Atlanta, Ga., November 1974.
Chertkov, B.A. "Mass Transfer Coefficients During Absorption of
SOp from Gases using Magnesium Sulfite and Bisulfite Solution,"
Khem. Prom. 7. 537-41 (1963).
107
-------
7. Chertkov, B.A. "Oxidation of Magnesium Sulfite and Bisulfite
during Extraction of SCL from Gases." J. App., USSR 33 (60)
2136-42 (1960).
8. Maxwell, M.A., and Koehler, G.R. "The Magnesia Slurry S02 Recovery
Process Operating Experience with a Large Prototype System,"
Presented at American Institute of Chem. Engineers 65th Annual
Mtg., New York, November, 1972.
'9. Quigley, C.P. and J.A. Burns, "Assessment of Prototype Operation
and Future Expansion Study - Magnesia Scrubbing, Mystic Generating
Station," Paper presented at EPA Flue Gas Desulfurization Symposium,
Atlanta, Georgia, November, 1973.
TOR
-------
Table 7.1 MAGNESIUM OXIDE PROCESS UNIT USAGE AND COST DATA
A. Chemicals & Utilities
Basis
Unit Cost
MgO: (based on 5% loss)
Coke:
Power (a) Absorbing
(b) S02 regeneration
Make-up Water
Fuel Oil
0.0322 lb/lbS02
0.0105 lb/lbS02
4,0 KW/M SCFM
0.10 KWhr/lbS02
0.2 gal/lb S02
0.037 gal/lbS00
$140/ton
$26/ton
$0.015/KWH
$0.30/M gal
$0.30/gal
B. Operating Labor &
Maintenance
Basis
Unit Cost
Labor (based on 365 day/yr)
S0_ Absorption
SO- Regeneration
Maintenance
Insurance and taxes
% man/shift
<75,000 SCFM
1 man/shift
>75,000 SCFM
2 man/shift
<10,000 lb/lbS02
2% man/shift
>10,000 lb/lbS02
1 man/day
5% TCI/yr
2 1/2% TCI/yr
$8/hr
Fixed Charges 13.12% TCI/yr
Based on Capital Recovery Factor using 10% interest over 15 year life.
109
-------
MAGNESIUM OXIDE SCRUBBING
TOTAL CAPITAL INVESTMENT COSTS
in
te
8
LU
2
w
LU
>
z
O-
<
o
o
S02 REMOVAL EFFICIENCY = 90%
SO,
$10 Million
NOTE: GAS COOLING & CONDITIONING
NOT INCLUDED.
1.0 MILLION
Costs: Midi:
100,000 SCFM
GAS FLOW RATE-SCFM
no
FIGURED
-------
MAGNESIUM OXIDE SCRUBBING
TOTAL ANNUAL DIRECT OPERATING COSTS
S02 RECOVERY EFFICIENCY = 90%
$1-0 Million
N°TE: GAS COOLING & CONDITIONING
NOT INCLUDED
4.0% SO,
2.0% SO.
SO,
0.5% S02
Costs: Mid 1974
100,000 SCFM
GAS FLOW RATE-SCFM
111
FIGURE 7.3
-------
MAGNESIUM OXIDE SCRUBBING
TOTAL CAPITAL INVESTMENT COSTS
SO2 REGENERATION SECTION
$1.0 Million
CO
fe
O
o
1-
z
LU
2
in
K.
<
o
O
Costs: Midi
10,000 LB/HR.
SULFUR DIOXIDE RATE
OF REGENERATION SYSTEM (Ibs/hr)
1)2
FIGURE]
-------
Table 7.2 MAGNESIUM OXIDE SCRUBBING
CAPITAL AND TOTAL ANNUAL COSTS
TOTAL CAPITAL INVESTMENT
Annual Cost
A. Direct Operating
1. MgO
2 . Coke
3 . Power
4. Fuel oil
5. Water
6 . Labor
7. Maintenance
8. Taxes & Insurance
TOTAL ANNUAL DIRECT
OPERATING COST
B. Annual Capitalized
Cost
C. Net Total
Annualized Cost*
$
$/SCFM
$/ Annual ton
S02 Removed
$/yr/ton
S02 Removed
$/yr/ton
SO- Removed
70,000
3,800,000
54
145
117,500
7,100
112,400
578,200
3,100
192,000
190,000
95,000
1,295,300
50
499,700
1,051,700
40
Gas Flow Ra
100,000
4,800,000
48
129
167,800
10,100
160,600
826,000
4,500
226,800
240,000
100,00
1,755,800
47
631,200
1,390,600
37
te - SCFM @1% S
200,000
7,250,000
36
97
335,600
20,200
321,100
1,651,900
8,900
262,000
362,500
181,300
3,143,500
42
953,400
2,356,000
32
°2
300,000
8,700,000
29
78
503,400
30,400
481,700
2,477,900
13,400
262,000
435,000
217,500
4,421,300
40
1,144,100
3,164,800
28
*Based on Corporate Tax Rate of 48%.
-------
Table 7.3. MAGNESIUM OXIDE SCRUBBING
TOTAL SYSTEM CAPITAL AND ANNUAL COSTS
: :-
A. TOTAL CAPITAL INVESTMENT
1. Gas Conditioning
2. S02 Absorption &
Recovery
3. Sulfuric Acid Plant
TOTAL
B. TOTAL ANNUAL
OPERATING COST
1. Gas Conditioning
2. S02 Absorption &
Recovery
3. Sulfuric Acid Plant
TOTAL
C. NET TOTAL
ANNUALIZED COSTS*
1. Gas Conditioning
2. S02 Absorption &
Recovery
TOTAL
$ /Annual ton
S02 Removed
$/yr/ton
SO 2 Removed
$/yr/ton
S0~ Removed
\
70,000
1,800,000
3,800,000
2,000,000
7,730,000
306
273,100
1,295,300
245,700
1,814,100
72
321,100
1,051,700
326,,800
1,699,600
67
\
Gas Flow R
100,000
2,275,000
4,800,000
2,520,000
9,800,000
272
370,600
1,755,800
295,700
2,422,100
67
419,100
1,390,600
404.600
2,214,300
61
ate - SCFM @1% !
200,000
3,550,000
7,250,000
3,950,000
15,000,000
208
613,500
3,143,500
474,500
4,231,500
59
672,200
2,356,000
639.700
3,667,900
51
j \J f\
300,000
4,650,000
8,700,000
5,400,000
18,750,000
173
843,300
4,421,300
619,700
5,884,300
54
740,000
3,164,800
859.500
4,764,300
44
-------
8.0 DIMETHYLANILINE/XYLIDINE PROCESS
8.1 PROCESS DESCRIPTION
In this process aromatic amines such as N,N-dimethylaniline (DMA)
or 2,3-dimethylaniline (xylidine) are used to absorb the SO,, from the
gas stream. In the case of xylidine, the absorbent is an aqueous
solution of xylidine, whereas the dimethylaniline is used essentially as
100 percent organic. The process is represented in Figure 8-1.! There
are three distinct steps which make up this process:
1) S02 Absorption ~ The cooled and cleaned flue gas is introduced
into the bottom of a special bubble-cap type of absorber column containing
three separate sections. The bottom section, where essentially all the
absorption of S02 takes place, uses approximately eight trays. The
second section, with usually only two or three trays, uses sodium carbonate
to remove the residual sulfur dioxide. The third section, with its nine
°r so trays and sulfuric acid feed, scrubs the flue gas of the DMA/xylidine
carried over from the bottom section.
Each of these units in the absorber tower is separate and distinct
iri operation and is provided with independent inlet and outlet nozzles
°r liquid flows. The liquid phases are kept separate for subsequent
treatment.
In the DMA system, as the S0? is absorbed by the anhydrous DMA, its
color changes from light yellow to deep ruby red.
Considerable heat is evolved during the absorption of sulfur dioxide
nd to optimize the absorption process, inter-tray coolers are provided
011 the bottom absorption section.
2) SQ2 Absorption — The S02~enriched DMA/xylidine stream is
from the absorber via an appropriate heat exchanger to a second
ulti-section bubble-cap type column. The S02-pregnant stream is
ntroduced into the middle or stripper section of this column, and as it
Passes down the four or five trays in this section, the S02 is liberated
the rising steam and S00 stream from the bottom or regenerator section.
Th
6 regenerator section with usually seven or eight trays, treats the
dilute sulfuric acid and sodium carbonate (sulfite and sulfate)
ams from the main absorber column to release dimethylaniline/xylidine
115
-------
vapor and S09 which passes up through the stripper section. This total
vapor load of S09 and dimethylaniline/xylidine then passes into the top
or rectifier section of the column where it is scrubbed with water. The
DMA/xylidine vapors in the presence of S02 are converted to the sulfite
form, dissolved in the water which overflows into the stripper section
beloW and joins the rich absorbent stream. The DMA/xylidine stream from
the bottom of the stripper section is transferred via the heat exchanger
to a separator where the DMA/xylidine separates and is returned to the
S02 absorption section.
The SCL stream leaving the rectifier is cooled and scrubbed with
water to remove the last traces of DMA/xylidine before passing to the
S0« utilization section.
A purge stream containing the sodium sulfate is removed from the
bottom of the regenerator section.
3) S02 Utilization — Although any SO, utilization plant can be
coupled to the DMA/xylidine Process, an SO, liquefaction plant is common,
and this system has been considered as being integral with this particular
SO, control process.
The cold S02 stream from the after scrubber of the regeneration
section is dried in a drying tower using 98 percent sulfuric acid,
compressed, cooled and run to storage.
8.2 PROCESS AND OPERATING CONSIDERATIONS
The process and operating considerations are similar for both the
anhydrous dimethylaniline and the aqueous xylidine processes although
there are some important differences.
Since the DMA system is essentially anhydrous, it is not particularly
susceptible to oxidation in the absorber section. The small amount of
DMA sulfate formed is readily converted to DMA and sodium sulfate during
the regeneration process, and the Na2S04 is removed as part of the
purge. However, DMA sulfate is also formed in the top section of the
absorber column where the DMA vapors are scrubbed with dilute sulfuric
acid. This sulfate is converted back to DMA and sodium sulfate in the
regeneration loop but the amount of sodium sulfate formed and purged is
related to the S02 concentration in the gas feed. As the SO, concentration
in the gas feed drops, the production of sodium sulfate per unit of S02
removed increases dramatically. At 5 percent SO,,, sodium sulfate format.**
116
-------
TO
INLET GAS
FRESH WATER
SEPARATOR WATER
CONDITIONED
GAS
ABSORBER
DIMETHYLAIMILINE/XYLIDINE PROCESS
FIGURE 8-1
-------
is on the order of AO Ibs/ton SCL recovered, but at 0.5 percent S02 in
the gas feed, the sulfate rate increases to 400 Ibs/ton SO- recovered.
At the lower concentrations of S0«, the solubility of S02 in DMA is
reduced. Pastnikov and Astashera^ have reported an equilibrium solubility
of S09 in DMA of 250 grams/liter for an S02 gas stream at 5 percent
against a value of 25 grams/liter for an S02 gas stream at 0.5 percent.
Hence at lower S0? concentrations, higher recirculation rates of DMA are
required and this tends to increase DMA losses, increase both Na_CO_
and H-SO, usage levels, and results in increased sodium sulfate generation.
On the other hand, the xylidine system with its higher solubility
capability at the low S0? concentrations is not penalized in this manner
when treating weak SO,, gas streams, but being an aqueous solution, it
does experience high oxidation rates in the absorber section itself as
the oxygen/S02 ratio increases. This situation requires special addi-
tional treatment of the absorbent stream with Na2CO~ to regenerate the
xylidine and the overall result is to increase significantly the load on
the regeneration section and the necessary sodium sulfate purge. This
higher stripper water bleed off to the regenerator is essential to
prevent the formation of solid reaction products within the system
itself through the reduced solubility of xylidine sulfate in water.
There is some uncertainty as to at what point one process begins to
demonstrate advantages over the other, although two authors with con-
siderable experience in this particular process have expressed the
opinion that the DMA system is competitive with the xylidine system down
3
to as low as 2 percent SO,,.
To enhance the S02 solubility capability of both process versions
but particularly the DMA system, intercoolers are required between'each
of the absorption section trays to remove the heat of absorption. An
advantage of these coolers is that they reduce the vapor pressure of the
DMA and, accordingly, reduce losses. This in turn reduces chemical
consumption and indirectly the formation of sodium sulfate.
8.3 PROCESS DEVELOPMENT STATUS
There are a number of DMA systems installed in both the U.S. and
overseas.
118
-------
The American Smelting and Refining Company (ASARCO) initially
applied the DMA process to the recovery of S0« from lead sinter machine •
gases at its Selby, California lead plant (now closed) in the late
1940' s. This plant recovered approximately 20 tons of liquid SC^/day
from gases with an SCL content ranging from 4-6 percent. ASARCO com-
pleted a second DMA plant at its Tacoma, Washington plant in 1973-74,
rated at 200 TPD liquid S09 processing copper converter gases of variable
^ *
SO content (3-10 percent) but averaging about 5 percent. This plant
is in operation today and is performing as expected.
ASARCO has also licensed construction of DMA plants as follows:
1) Phelps Dodge Corporatio, Ajo, Arizona — This plant was
nominally rated at 160 TPD liquid S02 and originally was to
treat mixed copper converter and reverberatory gases. Although
the plant was commissioned in 1973, operation has been only
spasmodic with numerous mechanical and equipment problems.
The plant has been modified and it is expected to start up at
reduced capacity on reverberatory gases only (1-3 percent S02)
during the second half of 1975.
2) Cities Service Company. Copperhill. Tennessee — Two plants
rated at 40 and 55 TPD liquid S02 were eventually brought into
operation. The gas feed uses a mixture of gases from iron and
copper roasters, reverberatories and converters with an S02
content of about 6 percent.
3) General Products Company. S.A. Mexico City — Plant rated at
5 TPD liquid S02 treating gases containing 16 percent S02-
4) Real Compania Asturiana de Mines. Torrelavega, Spain — Plant
capacity 165 TPD liquid S02.
It is apparent that the DMA process operating on S02 streams with
concentrations above 4 or 5 percent is a proven commercial process. Its
effectiveness, either in the DMA or xylidine variation, on weaker S02
streams is much less certain. The Ajo installation is yet to demonstrate
lt:8 capability on low S02 content streams, and the literature references
to commercial application of the xylidine (or sulphidine) process in
Europe are dated and discuss small installations prior to 1945. One
in the Norddeutsche Affinerie in Hamburg during the 1930' s
rePortedly used the xylidine process to treat copper converter gases
c°ntaining 0.5-8.0 percent S02 (average 3.6 percent).
119
-------
8.4 PROCESS DESULFURIZATION EFFICIENCY
This process appears to be capable of very high S09 recovery
56
efficiences. A number of references ' have indicated efficiency levels
of above 98 percent but in each case, the process was applied on streams
-------
been necessary to install mist precipitators after the cooling and
humidifying towers located after the dry electrostatic precipitator on
the reverberatory gas system.
8.7 PARTICULATE REMOVAL CAPABILITY
The DMA/xylidine absorber has minimal particulate removal capability
^specially on gas streams which have already received an extensive
precleaning sequence.
8.8 PROCESS ENERGY REQUIREMENTS
This process is not an energy intensive process. Blower horsepower
requirements might be expected to be somewhat higher than for other
scrubber systems because of the higher pressure drop associated with the
special absorber column. Total power requirements (Kwh) per ton of SO™
recovered will be inversely related to the concentration of SO™ in the
gas feed, with low strength streams incurring the high power values.
Steam requirements will be greater with the xylidine approach than with
DMA, but values of between 1 to 1.5 Ibs/ton S02 appear likely.
8-9 RETROFITTING REQUIREMENTS
The DMA/xylidine process with its gas conditioning requirements and
the necessity to locate the special absorber column and its associated
intercoolers and secondary loops in proximity to the existing flue
ductwork will pose retrofitting problems in any old or congested smelter.
lt: is possible to locate the stripper section and S02 liquefaction
Section away from the absorption area, but based on the experiences at
Tacoma and Ajo, DMA plants require at least 40-50,000 sq. ft. of total
lnstallation area.
8-10 PROCESS COSTS
The installations of DMA units at Tacoma and Ajo within the past 4
°r So years have provided overall installed cost figures for two in-
flations similar in rated SO- capacity—200 TPD and 160 TPD, respectively-
and both handling gas streams of quite variable flow and SO- concentrations.
As is commonly the case with such overall cost figures, and without the
ei*efit of detailed work scope breakdowns, the two costs are difficult
to evaluate. Some order of magnitude estimates have also been offered
j Q
n c°rrespondence between ASARCO representatives and EPA. This data
as been used after appropriate adjustment and escalation together with
121
-------
the background information on both the DMA and xylidine processes to
develop the capital and operating costs presented in Figures 8-2 and 8-
3, respectively. The emphasis has been on a process treating low SC^
content gases, and the operating costs have been extrapolated to reflect
the impact of weak SCL streams on the cost structure. A cost study on
9
the DMA Process prepared by Allied Chemical was also reviewed in determining
operating unit usage values.
It should be noted that with this process, although the capital
costs tend to show a fairly narrow increasing spread of costs with SO-
concentration, the highest operating costs are associated with the
weakest SO- gas streams. As the gas stream S02 concentration increases,
the operating costs are reduced, but at 4 percent SCL content the annual
operating costs appear to increase. Obviously, this response reflects
the way the costs, both capital and operating, have been structured.
For this process only, the SO- utilization section—i.e., compression
and liquefaction—has been included in the basic process cost. Also
because of the integrated nature of the S02 regeneration circuit and the
absorption loop, no attempt has been made to provide a separate SO-
regeneration cost curve as was done for the other candidate control
processes.
Tables 8.2 and 8.3, respectively, provide capital and direct operating
cost breakdowns for this process applied to a range of gas flows con-
taining 1 percent SO,, and a total system cost including gas conditioning
for the same range of gas flows and SO- concentration.
8.11 ADVANTAGES AND DISADVANTAGES
Advantages of this process are:
1) Based on the experience with higher SO,, streams, this system
has few significant operating problems if the gases are
adequately cleaned and conditioned.
2) The process achieves a high level of S02 removal efficiency
over a wide range of SO- concentrations':
3) On higher S02 content gas streams, the DMA system provides
good control of oxidation and is capable of treating higher
oxygen level streams without penalty.
122
-------
The disadvantages of the process are:
1) The xylidine alternative for treating low S02 content streams
is a relatively undemonstrated process.
2) Oxidation rates in the xylidine system treating low SO-
content gases are relatively high and increase with decreasing
S00 concentration. Control becomes more critical.
3) Chemical consumption and steam requirements increase as
the S0~ concentration drops.
8.12 REFERENCES
1. Office of Water and Air Programs, EPA, Background Information for
New Source Performance Standards: Primary Copper, Zinc, and Lead
Smelters, February, 1974, pp. 4-65 to 4-74.
2. Pastinikov and A.A. Astasheve. J. Chem. Ind. (USSR) 17(3): 14-19
(1940.)
3. Fleming, E.P. and T.C. Fitt, "Liquid Sulfur Dioxide from Waste
Smelter Gases," Industrial and Engineering Chemistry, November
1950, pp. 2253-2258.
4. Henderson, J.M. and J.B. Pfeiffer, "Liquid S02 from Copper Smelter
Gases - The ASARCO DMA Process," presented at the 78th National
Meeting, American Institute of Chemical Engineers, Salt Lake City,
Utah, August 1974.
5- Meisel, G.M., "Sulfur Recovery," Journal of Metals, May 1972,
pp. 31-39.
6- Weidmann, H. and G. Rosener, Industrial and Engineering Chemistry,
New Edition, March 1936, p. 105.
7- Kohl, A.L. and F.C.'niesanfeld, "Gas Purification," McGraw-Hill,
New York, 1960, pp. 197-209.
8- Henderson, J.M., ASARCO. Private communication to Jack R. Farmer,
EPA, March 15, 1973.
9- "Applicability of Reduction to Sulfur Techniques to the Development
of New Processes for Removing S0? from Flue Gases, Vol. II,
Allied Chemical Corporation, Morristown, New Jersey, November 1970,
NTIS PB 198-408.
123
-------
Table 8.1. DMA/XYLIDINE PROCESS
UNIT USAGE AND COST DATA
A. Chemicals & Utilities
Basis
Unit Cost
DMA/Xylidine
Sulfuric Acid
Soda Ash
Electric Power
Cooling Water
Steam
(a)
(b)
002
%so2
100
Ib/MSCFM
j).Q34
100
0.03
Ib/MSCFM
100
Ib/MSCFM
8.0 KW/MSCFM
0.02 KWh/lbS0
300
gal/MSCFM
1.5 lbs/lbS0
$ 0.50/lb
$ 48/ton
$ 52/ton
$ 0.015/KWh
$ 0.1/Mgal
$ 1.25/Mgal
B. Operating Labor &
Maintenance
Labor
Maintenance
Insurance & Taxes
2 man/shift
5% TCI/yr
TCI/yr
$ 8/hr
C. Fixed Charges 13.15% TCI/yr
Based on Capital Recovery Factor using 10% interest over 15 year
124
-------
DMA/XYLIDINE SCRUBBING
TOTAL CAPITAL INVESTMENT COSTS
S02 REMOVAL EFFICIENCY 95%
$10.0 Million
4.0% S02
2.0% SO2
1.5% SO2
1.0% S02
0.5% S02
NOTEE.
GAS COOLING & CONDITIONING
NOT INCLUDED.
Costs: Mid 1974
100,000 SCFM
GAS FLOW RATE-SCFM
FIGURE 8-2
-------
DMA/XYLIDINE SCRUBBING
TOTAL ANNUAL DIRECT OPERATING COSTS
SO2 REMOVAL EFFICIENCY 95%
0.5% SO2
$10.0 Million
NOTE: GAS COOLING & CONDITIONING
NOT INCLUDED.
$1.0 Million
4.0% SO
'2
.0-2.0% S02
Costs: Mid 197i
100,000 SCFM
GAS FLOW RATE-SCFM
126
FIGURE
-------
Table 8.2, jyXMEIHYLAfflLIifE/XYLIZHfrE PROCESS
CAPITAL AND TOTAL ANNUAL OPERATING COSTS
TOTAL CAPITAL INVESTMENT
Annual Cost
A. Direct Operating
1. Power
2. Water
3. Steam
4. Soda Ash
5. Sulfuric Acid
6. DMA/Xyline
7 . Labor
8, Maintenance
9. Taxes & Insurance
TOTAL ANNUAL DIRECT
OPERATING COST
B. Annual Capitalized
Cost
C. Net Total
Annualized Cost*
$
$/SCFM
$/ Annual Ton
SO-Removed
$/yr/ton
SO-Removed
$/yr/ton
SO Removed
Gas Flow Rate - SCFM @1% S02
70,000
8,000,000
114
291
85,000
102 , 800
102,800
267,300
279,700
342,700
140,200
400,000
200,000
1,920,600
70
1,052,000
1,794,700
65
100,000
10,230,000
102
261
121,500
146,900
146,900
381,900
399,500
489 , 600
140,200
511,500
255,800
2,593,800
66
1,345,000
2,366,000
60
200,000
16,200,000
81
206
243,000
293,800
293,800
763,800
799,000
979,200
140,200
810,000
405,000
4,727,800
60
2,130,300
4,070,400
52
300,000
21,560,000
72
183
364,500
440,600
440,600
1,145,700
1,198,500
1,468,800
140,200
1,078,000
539,000
6,815,900
58
2,835,000
5,689,400
48
to
Based on Corporate Tax Rate of 48%
-------
Table 8.3. DIMETHYLANALINE/XYLIDINE PROCESS
TOTAL SYSTEM CAPITAL AND ANNUAL OPERATING COSTS
A. TOTAL CAPITAL INVESTMENT
1. Gas Conditioning
2. S02 Absorption and
Recovery
TOTAL
B. TOTAL ANNUAL
OPERATING COSTS
1. Gas Conditioning
2. S02 Absorption and
Recovery
TOTAL
C. NET TOTAL
ANNUALIZED COSTS*
1. Gas Conditioning
2. S02 Absorption and
Recovery
TOTAL
$/ Annual ton
SO. Removed
$/yr/ton
S0? Removed
$/yr/ton
SO „ Removed
Gas Flow Rate - SCFM @1% S02
70,000
1,800,000
8,000,000
9,800,000
356
273,100
1,920,600
2,193,700
80
321,100
1,794,700
2,115,800
77
100,000
2,275,000
10,230,000
12,505,000
319
370,600
2,593,800
2,964,400
76
419,100
2,366,000
2,785,100
71
200,000
3,550,000
16,200,000
19,750,000
252
613,500
4,727,800
5,341,300
68
672,200
4,070,400
4,742,600
60
300,000
4,650,000
21,560,000
26,210,000
226
843,300
6,815,900
7,659,200
65
740,000
5,689,000
6,429,000
55
to
00
-------
9.0 CITRATE PROCESSS
9.1 PROCESS DESCRIPTION
The citrate process as it is presently defined uses an aqueous
solution of citric acid, sodium citrate and sodium thiosulfate to scrub
the S02-containing gas; the S02~containing absorbent is treated with H2S
to precipitate elemental sulfur; part of the sulfur produced is recycled
for production of H,,S while the balance is removed as product.
The overall process is comprised of the following steps or sections
and is depicted in Figure 9-1 which follows the flowsheet provided by
the Smelter Control Research Association in their report to the Bureau
of Mines.1
1) Sulfur Dioxide Absorption ~ The input S02~containing gas,
cooled to between 108-125°F (42-52°C) and cleaned of particulate material
in an appropriate gas conditioning system, is contacted countercurrently
in either a packed column or an impingement plate column with the aqueous
citrate solution. The treated gas passes through an entrainment separator
before discharge to the stack. The S02-loaded citrate solution is
Pumped to the regeneration section.
2) Regeneration ~ The S02-rich solution from the absorber feeds
a cascade of three agitated regenerator vessels where it is contacted
countercurrently with an H2S stream to precipitate elemental sulfur.
The reaction is moderately exothermic and is maintained at a temperature
°* around 150°F (65.56°C). If the H2S is generated by the interaction
of methane and sulfur, the gas stream is actually a mixture of H2S and
c°2 and the first vessel is vented back to the S02 absorber or to an
aPPropriate incinerator to oxidize the excess H2S. The sulfur slurry
flows to a conditioning tank where kerosene (or SAE 10 motor oil) is
introduced to agglomerate and float the sulfur. This material containing
about 50 percent solution is skimmed and transferred to a melter where
th-e sulfur is melted at about 275°F (135 °C). Molten sulfur and citrate
s°lution pass into a closed settler tank under a pressure of 35 psi.
H°lten sulfur is tapped from the bottom for casting into blocks or
Pumping to the H2S generation section. Citrate solution and oil are
Wlthdrawn from the top of the settler through a knockout pot, filter,
C°°ler and decanter for re-use. The regenerated citrate solution from
the settler passes through a settling tank and after filtration is
eturned to the absorption tower.
129
-------
3) ELS Generation — The filtered liquid sulfur is preheated,
vaporized and superheated to 122°F (649°C). Part of the steam and
natural gas required for H2S generation is also preheated and mixed with
part of the sulfur vapor before entering the catalytic H~S reactor. The
remaining sulfur vapor is blended with an unpreheated steam-natural gas
mixture and is injected into the catalytic bed to maintain near isothermal
bed temperature. The heat of the reaction products, H2S and C02> is
used to preheat the sulfur feed and for steam generation before final
cooling to about 140°F (60°C) and transfer to the regeneration section.
4) Sulfate Purge — A slipstream from the regenerated citrate
solution to the SO,, absorber is treated in a crystallization unit where
the solution is cooled to about 40°F (5°C) to crystallize out Glauber's
salt or sodium sulfate decahydrate. The recovered mother liquor returns
to the regenerated citrate solution loop and the sodium sulfate crystals
are disposed of.
It should be noted that the Bureau of Mines is currently working on
a process option to the normal process whereby the SCL-enriched absorbent
is steam-stripped to release the S02 rather than proceeding through the
H2S reaction step and recovery of elemental sulfur. Steam consumption
rates are at present high; rates of 10 Ib steam/lb SO, are required when
1 z
absorbing from a 0.5 percent S02 inlet gas.
9.2 PROCESS AND OPERATING CONSIDERATIONS
The citrate process uses the citrate ion, provided by the presence
of citric acid and sodium citrate in an aqueous solution, as a buffering
agent to increase the solubility of S02 in water. Under usual conditions,
the absorption of S02 in water is PH dependent, and as hydrogen ions are-
released by the dissociation of the sulfurous acid, it becomes self-
limiting. The equilibrium condition may be expressed as:
S02 + H20 2 HSO~ + H+.
The buffering action of the various citrate species removes the hydrogen
ions and allows substantially higher S02 loadings to be attained by the
absorbent. Thus the solubility of S02 in an aqueous solution is a
130
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SEPARATOR
WATER
SULFUR
VAPORIZER
r
I
I GAS
I CONDITIONING I
SECTION I
1 1
•^
\ *•
1
1
A
* i/ * ft
—
Y
1
1
1
1
INLET GAS FROM
DRY GAS CLEANING
SECTION
OIL
RECOV-
ERY
DRUM
S02 ABSORBER
KEROSENE
OIL
SEPARATOR
SULFUR REGENERATORS
SULFUR SKIMMER
CRYSTALLIZATION
UNIT
CONDITIONER
TANK
LIQUID
SULFUR
SETTLER
TO DISPOSAL
CITRIC ACID STEAM
SULFUR
STORAGE
PIT
ABSORBER
FEED TANK
FIGURE 9.1
_NATURAL
GAS
STORAGE TANK
CITRATE PROCESS
-------
function of the partial pressure of S02 in the gas phase, the hydrogen
ion concentration and the ionization constants of sulfurous acid. The
important relationships in the citrate process can therefore be identified
as:
1) the solution pH
2) the citrate concentration
3) the concentration of the SCL in the gas feed
4) the temperature of the gas feed.
Sulfur dioxide absorption is favored by increased pH and buffer
content of the solution, increased S02 content of the gas feed, and
decreased temperature. Under operating plant conditions, the citrate
concentration would be established at the lowest level compatible with
the SO- content of the gas and requisite solution flows, and the operating
temperature selected would balance the advantages of higher SCL loading
of the absorbent at lower temperatures against the increased cost of gas
cooling. The absorbent recirculation rate used for a specified S02
removal efficiency will depend on the actual citrate concentration
selected to satisfy the S02 content of the gas input. In the Bureau of
Mines Bunker Hill pilot plant program, the absorbent flow rate on a gas
stream of 1000 SCFM containing 0.5 percent S02 was about 10 gal/minute j
for a 0.5 M citrate solution to give an SO,, loading of about 10 grams/litef'
The pH limit is controlled by the regeneration step rather than the
absorption step, but the test work to date on the above gas stream has
been conducted in the pH range of 4.0 to 4.6 using 0.5 M citrate solution
with a molar ratio of NaOH to citric acid of 2. Operating temperatures
have been in the range of 108 to 149°F (42 to 65°C).
In practice, when regenerated absorbent is returned to the absorptio11
loop, thiosulfate is introduced into the system. The thiosulfate ion
complexes with the HSO~ and H formed in the absorption process, and
thereby sharply depresses oxidation of the HSO~. SO- absorption itself
is also aided by the presence of the thiosulfate. The presence of
thiosulfate in the absorbent solution is therefore an essential element
of the citrate process.
The chemistry of the citrate process is discussed in some detail by
o
a Bureau of Mines publication and a paper presented at the American
Chemical Society National Meeting in 1974.4 The overall stoichiometry
of the regeneration reaction is the same as the gas-phase Glaus reactio0'
132
-------
but the chemistry is more complex than that represented by the usual
reaction:
S02 + 2H2 + 3S + 2H20.
At pH 4.0 to 4.5, the usual PH of the citrate absorbent solution,
thiosulfate, trithionate, tetrathionate, and polythionates are all
believed to be formed in the regeneration step, with further reactions
with H2S yielding elemental sulfur. The discussion of the regeneration
chemistry provided in the paper by representatives of the McKee-Peabody-
Pfizer group indicates that the reaction of H2S with thiosulfate is the
rate determining step in regeneration and that by allowing the thiosulfate
concentration to build up in the system, the reaction rate is increased
and smaller reactors can be used. They also note that an important
feature of the citrate process is the capability of the buffering capacity
°f the citrate content and the thiosulfate content to minimize short-
term overloads of either S02 or H2S.
As with all closed Ipop systems, the cumulative buildup of oxidation
Products, particulates, and other soluble impurities contributed by the
fflakeup water and process chemicals must be controlled by purging a part
°f the recirculated absorbing solution. The addition of soda ash,
2^3' to *-he recirculation loop of the absorbing solution produces
2^4 from the oxidation sulfate ion product which is readily removed
after crystallation and filtration. This treatment will also tend to
remove particulate matter by occlusion. A further treatment may be
necessary to remove soluble impurities but the frequency and extent of
this treatment will depend upon the characteristics of the makeup water
and purity levels of the process chemicals used.
Although the process is being developed to yield elemental sulfur,
Variations of the process are possible. The S02~loaded citrate solution
Can be steam stripped, and depending on the level of solution loading,
about two-thirds of the S00 can be recovered using 3 to 5 Ibs of steam
D 3
per lb S02. The residual SCL in the solution would be precipitated in
116 normal manner by HLS for H2S generation with the regenerated citrate
°lution being returned to the scrubbing loop. This is an alternative
133
-------
approach to complete stream stripping which requires almost twice the
steam consumption. The stripped SO™, after condensation of the water
vapor, can be used as feed with approximate air dilution to a sulfuric
acid or used to enrich a lean SO™ gas stream for feed to a sulfuric acid
plant. These options provide considerable flexibility in both sizing
and mode of application to optimize around the local conditions which
may exist at any specific smelter.
9.3 PROCESS DEVELOPMENT STATUS
The citrate process is a second generation SO™ control process and
is still in the pilot plant developmental stage.
A small pilot plant to process up to 300 cfm of reverberatory
furnace gas was placed into operation jointly by the Bureau of Mines and
Magma Copper Company in November 1970 at the San Manuel, Arizona Copper
Smelter. Although operation was intermittent over a a six month period
due to equipment breakdowns and operational problems, the capability of
the SO™ absorption and regeneration system to remove 93 to 99 percent of ,
the SO™ from the 1.1 to 1.7 percent SO™ content smelter gas was demonstrated-
In 1969, Arthur G. McKee and Company, as part of a study for the
U.S. Department of Health, Education and Welfare on SO™ emissions from
non-ferrous smelters, made a broad survey and analysis of the technology
applicable to SO™ emission abatement. From this survey, they established
criteria for an ideal SO™ control process, criteria which appeared to be
met by the citrate process. In September 1972, McKee, Peabody Engineering
and Pfizer announced a joint project to install a 2000 SCFM pilot plant
unit on a coal-fired steam boiler at Pfizer*s Terre Haute, Indiana plant -
to generate hard engineering and economic data. Operation began in June
1973.
The Bureau of Mines development program has also continued with
laboratory work and a pilot plant installation at the Bunker Hill Company
2 5
lead smelter in Kellogg, Idaho to treat 1000 SCFM of 0.5 percent SO™.
Phase I of this program, covering the SO™ absorption and sulfur recovery
sections, was completed in January 1974 and operation began in February
1974 on a clean diluted 0.5 percent S02 stream from a sintering furnace.
Construction of Phase II, the H™S generation plant to produce 76 to 78
percent H™S from natural gas and steam, was completed in September 1974.
134
-------
Phase III, the gas cooling and cleaning section of the dirty sinter
plant tail gas containing 0.3 to 0.9 percent S02, was scheduled for
completion early in 1975.
Although the programs have confirmed the system chemistry, expected
SO- removal efficiency, and the essential operating relationships of the
citrate process, there is still a requirement to demonstrate the full
commercial potential of the process under extended plant operating
2
situations and conditions. The Bureau of Mines has recently announced
that plans are underway via a cooperative and cost-sharing arrangment
between the Bureau of Mines, EPA and interested industrial firms to
provide one or more large scale demonstration plants at power plants or
steam generating facilities burning high sulfur coal or oil.
The short continuous run times common in pilot plant installations
do not always provide adequate data on the operating losses which might
be experienced in full scale commercial operations. The use of kerosene
and the losses per ton of sulfur recovered presently estimated by the
Bureau of Mines constitute a major part of the total raw material cost.
However, laboratory scale tests have indicated that the use of SAE 10
motor oil provide results equivalent to those produced by kerosene with
2
an oil consumption one-fourth that of kerosene. The McKee-Peabody-
Pfizer version of the process uses no hydrocarbon addition but affects
sulfur separation based on a flotation principle. The cost advantages
of this approach are obvious but it is not known if there are offsetting
difficulties. Very little experience has been gained in pilot plants by
using on-site generated H?S. A reliable source of H2S is mandatory to
insure uninterrupted operation of the citrate process and while commercial
quantities of H2S can be obtained from several alternative sources, the
locations of non-ferrous smelters suggest that economic factors would
tend to emphasize on-site generation using some form of sulfur, steam
and a reducing agent. A recent study reviews the potential of leaching
°f neutral roasted copper concentrations with HCl and the associated
release of H-S. Such an operation would mesh well with an on-site
citrate process. The continuing program at the Bunker Hill pilot plant
Deludes demonstration of the sulfur-methane-steam reaction for H2S
generation to provide both engineering data and information on the
influence of impure, on-site generated 76 to 78 percent H2S on sulfur
Precipitation.
135
-------
The McKee-Peabody-Pf izer pilot plant program at Terre Haute has
indicated that the level of gas cleaning required for the citrate process
may not be as demanding as that originally considered to be necessary
and that particulate material can be readily removed from the regeneration
loop. As noted above, Phase III of the Bureau of Mines Bunker Hill
pilot plant which is scheduled for operation in 1975 will focus on
defining gas cleaning requirements and the effect on process capabilities.
9.4 PROCESS DESULFURIZATION EFFICIENCY
2 3
Pilot plant programs conducted by both the Bureau of Mines ' and
4
the McKee-Peabody-Pfizer group have consistently obtained S02 removal
efficiencies between 93 and 99 percent. SCL content of the inlet gas
streams has ranged from 0.1 to 0.2 percent for the flue gases from a
coal-fired steam generating boiler to 1.0 to 1.7 percent for a copper
reverberatory furnace stream. It is expected that any full scale in-
stallation, using the process as developed, could achieve an overall S02
removal efficiency of at least 95 to 96 percent.
9.5 SULFITE OXIDATION
The citrate process, with its various citrate species providing a
buffering action, increases the solubility of SO,., in water with formation
of the bisulfite ion:
Oxygen is universally present in smelter and combustion process flue
gases together with some H2SO, pickup from incomplete mist removal, and
during the absorption step, some oxidation of the .bisulfite occurs:
HSO~ + 1/2 00 -> HSOT * H+ + SO?.
_) Z q 4
Some oxidation may also occur in the solution present during melting of
the sulfur:
3HS03 + 250^ + H+ + S + H20.
136
-------
Certain metal ions such as copper and iron act as catalysts for the
reaction. Results from some laboratory runs reported by McKee-Peabody-
4
Pfizer indicate that the oxidation rate increases with pH, a change
from pH 4 to pH 5 approximately doubling the rate. The results also
indicate that the addition of thiosulfate reduces the oxidation rate.
The formation of thiosulfate during the regeneration step and the presence
of these ions in the absorption system thus tend to inhibit the oxidation
process. However, some decomposition of the thiosulfate does take place
during the melting of the precipitated sulfur, and this reaction produces
additional sulfate. The quantity of sulfate formed is dependent on the
concentration of thiosulfate, temperature, and duration of heating; but
on the basis of laboratory tests, the Bureau of Mines has estimated that
under expected operation conditions, the sulfate formation during melting
3
would be about 3 Ibs/ton sulfur melted.
During the operation of the Phase I section of the Bunker Hill
pilot plant with a feed gas containing about 25 percent oxygen, the rate
- 2
of oxidation of S02 to SO, was determined to be about 1.3 percent or
approximately 53 Ib/ton of sulfur melted. The thiosulfate concentration
in this period ranged between 20 to 40 gms/liter, a much lower concentration
than would normally be expected because of the large solution losses and
fresh citrate solution makeup required by plant upsets.
While the overall rate of sulfate formation from all sources is
small, the effect is cumulative and requires that sulfate be purged from
the system.
9.6 FEED GAS PRETREATMENT
To prevent excessive evaporation of the citrate solution in the
absorption tower, the feed gas must be cooled and saturated with water
vapor in a cooling/humidification system typically described in Appendix
A. Because S02 absorption by the citrate solution is favored by lower
temperatures, the selection of a preferred operating temperature will
represent a trade-off between the higher S02 loadings of the citrate
solution associated with lower temperatures and the effect on equipment
and piping costs and the cost of the gas cooling section itself. Pilot
Plant work at the San Manuel location was done at temperatures between
iOS to 125°F (42 to, 528C) but higher temperatures between 113 and 149°F
(45 to 65°C) were used in the Phase I test program at the Bunker Hill
137
-------
smelter to correspond with temperatures to which the lead sintering
2
furnace tail gas would be cooled in Phase III of the program.
The gas cooling and conditioning system described in Appendix A
would provide acceptable treatment for the process as its requirements
are now defined. The particulate loading remaining in the gas stream
after dry gas cleaning in the smelter itself would be further reduced by
both the quenching step and the electrostatic mist precipitator itself.
The proposed testing program at the Bunker Hill smelter under Phase
III will provide more detailed information on the overall degree of feed
gas pretreatment required by the citrate process when handling hot
smelter-generated gas streams.
9.7 PARTICULATE REMOVAL CAPABILITY
As with other S09 control processes, the S09 absorber in the citrate
£* £
process is not a high pressure drop device, and with a gas feed that has
received prior mechanical and electrostatic cleaning followed by water
quenching and mist electrostatic precipitation, the particulate capture
on smelter streams will be small.
The use of packed columns in the Bureau of Mines development program
will tend to provide somewhat better collection than a sieve tray or
impingement plate column, but with the particle size distribution pre-
dominately under 1 to 2 ym, the mass removal overall is negligible.
The process itself will tend to maintain any captured particulate
material at an equilibrium level since the sulfur precipitation step by
occlusion will continually remove such material from the closed loop of
the system.
9.8 PROCESS ENERGY REQUIREMENTS
The direct energy requirements in terms of electric power of the
citrate process are relatively modest. Power requirements for gas
movement though the conditioning and SO™ absorption sections will be
similar to those for other desulfurization processes, but lower absorbed
liquid recirculation rates will reduce pumping power costs accordingly-
Power costs in the regeneration and recovery sections are expected to be
around 0.075 KW/lb S02 recovered.
138
-------
The use of natural gas both as fuel and as process raw material
stock in the production of H^S constitutes a major part of the energy
demands of this process. Approximately 4 cubic feet of natural gas are
required for every pound of SC^ removed. For a gas flow of 100,000 SCFM
containing 1 percent S02> the annual cost of natural gas represents
approximately 27 percent of the total annualized operating costs.
However, where local conditions favor steam stripping of the SO- for
output to a sulfuric acid plant, this energy cost would be substantially
reduced.
9.9 RETROFITTING REQUIREMENTS
The regeneration and sulfur recovery sections can be located at
some distance from the gas conditioning S02 absorption sections and
existing plant configurations should not pose any undue difficulties.
Both the gas conditioning and S02 absorption sections must be
retrofitted to existing ductwork and within the structural and space
limitations of any existing plant. The age of the plant will exercise a
significant influence on the degree of difficulty and the associated
cost picture. Space requirements are estimated as follows:
Gas Conditioning and SO., Absorption 8-12,000 sq. ft.
(including neutralization and
cooling tower)
Regeneration and sulfur recovery 15-20,000 sq. ft.
9-10 PROCESS COSTS
Published cost data are limited and are based on projections from
the small pilot plant programs handled to date. Over the past several
years, the Bureau of Mines and Arthur G. McKee and Company have provided
estimates for both non-ferrous applications and power plants, ' ' and
the Lummus Company prepared for the Smelter Control Research Association,
^c.,1 an engineering estimate for a copper reverberatory furnace gas.
these essentially order-of-magnitude costs have been reviewed and have
p*°vided the basis for developing capital costs for the citrate process
te8eneration section on the basis of SO,, handled in Ibs/hr. The costs
f°r-various S02 rates have been combined with the capital cost of the
*elated S02 absorption section developed in Appendix II to provide the
Pa*ametric capital cost relationships detailed in Figure 9-2.
139
-------
Total annual operating costs for the citrate process based on gas
flow and concentration of S02 are provided in Figure 9-3.
For comparative purposes, Table 9-2 provides a capital and operating
cost breakdown for this process applied to a range of gas flows containing
1 percent SO^- Table 9-3 provides a total system cost, including gas
cooling and conditioning, and allows direct comparison with similar
system costs for the other SO,, control processes under evaluation.
9.11 ADVANTAGES AND DISADVANTAGES
Advantages of the process are:
1) The process achieves a high level of S0~ removal efficiency
over a wide range of SO- concentrations.
2) There are no potential scaling or plugging problems in the
absorption section.
3) The process is inherently stable and can accommodate variable
SO^ loadings without upset in the KLS regeneration section.
4) There is a minimum effect from oxidation with the citrate
absorbent solution providing an oxidation-inhibiting effect.
5) The process provides direct recovery to elemental sulfur with
an option of generating S02 for acid plant feed.
The disadvantages of the process are:
1) The process has been demonstrated only at the pilot plant
level and full scale commercial application is still several
years off.
2) The requirement for natural gas in the H2S generation section
for the elemental sulfur option is fairly high and constitutes
a significant part of the operating charges—25-30
percent of the total annual direct operating costs.
9.12 REFERENCES
1. Smelter Control Research Association, Inc. "Report to U.S. Bureau
of Mines on Engineering Evaluation of Possible High Efficiency
Soluble Scrubbing Systems for the Removal of S09 from Copper Smelter
Reverberatory Furnace Flue Gases," B.O.M. Contract No. S0133044,
March 1974.
2. W.A. McKinney, et al. "Pilot Plant Testing of the Citrate Process
for SCL Emission Control," Presented at Flue Gas Desulfurization
Symposium, Atlanta, Georgia, November 4-7, 1974.
3. J. B. Rosenbaum, W.A. McKinney, et al. "Sulfur Dioxide Emission
Control by H2S Reaction in Aqueous Solution, the Citrate Process,"
Bureau of Mines Report of Investigations RI 7774, 1973.
140
-------
4. L. Korosy, H. L. Gewanter, F. S. Chambers and S. Vasan. "Chemistry
of SO- Absorption and Conversion to Sulfur by the Citrate Process".
Paper presented at the Symposium on Sulfur Removal and Recovery
from Industrial Sources, 167th American Chem. Society National Mtg.,
Los Angeles, California, April, 1974.
5. J.B. Rosenbautn, D.R. George, and L. Crocker. The Citrate Process
for Removing SCL and Recovering Sulfur from Waste Gases. Presented
at A.I.M.E. Environmental Quality Conference, Washington, B.C.,
June 7-9, 1971.
6.
C.A. Rohrmann and H.T. Pullman. "Control of SCL Emissions from
Copper Smelters," Vol. II, Hydrogen Sulfide Production from Copper
Concentrates," Battelle Pacific Northwest Laboratories, EPA-650/2-
74-085-5, September, 1974.
7. Private communication with Frank S. Chambers, Assistant Director,
Consulting and Developing, Arthur G. McKee & Co. , Cleveland, Ohio.
8. F.S. Chambers, L. Korosy and A. Scaleem. "The Citrate Process to
Convert S09 to Elemental Sulfur". Paper presented at Industrial
Fuel Conference Purdue University, October, 1973.
9. Private communication with W.I. Nissen, Bureau of Mines, Salt Lake
City, Utah.
141
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Table 9.1. CITRATE PROCESS UNIT USAGE AND COST DATA
A. Chemicals & Utilities
Citric Acid
Sodium Carbonate
Kerosene
Natural Gas
Catalyst
Power a) S02 Absorption
b) S02 Regeneration
Water a) Process
b) Cooling
B. Operating Labor &
Maintenance
Labor: Absorption
S0? Regeneration
Maintenance:
Taxes and Insurace:
Basis
0.003 lb/lbS02
0.009 lb/lbS02
0.0036 gal/lbS02
3.6 CF/lbS02
0.00042 lb/lbS02
4 KW/M SCFM
0.075 KWhr/lbS02
1.5 gal/lbS02
5.0 gal/lbS02
% man/shift
<75,000 SCFM
3/4 man/shift
>75,000 SCFM
2h man/shift
<10,000 Ib/yr
31/4 man/ shi ft
>10,000 Ib/hr
4.0% TCI/year
2.5% TCI/year
Unit Cost
—~
$0.50/lb
$52/ton
$0.50/gal
$1.25/MCF
$0.15/lb
$O.Q15/KWH
$0.30/M gal
$0.10/M gal
$8/hr
^
-*^^
C. Fixed Charges 13.15% TCI/year
Based on Capital Recovery Factor using 10% interest over 15 year life.
142
-------
CITRATE PROCESS
TOTAL CAPITAL INVESTMENT COSTS
S02 REMOVAL EFFICIENCY = 95%
SO,
$10 Million
°TE: GAS COOLING & CONDITIONING
NOT INCLUDED.
0.5% S02
Costs: Mid 1974
100,000 SCFM
GAS FLOW RATE-SCFM
I VI
FIGURE 9.2
-------
CITRATE PROCESS
TOTAL ANNUAL DIRECT OPERATING COSTS
S02 REMOVAL Ef FICIENCY - 95%
4,0% SO-
V)
8
IT
111
O.
O
Ill
DC
D
-J
13
2.0% SO.
$1.0 Million
<
O
WOTE: GAS COOLING & CQNDITIQNJNG
NOT INCLUDED.
Costs:
100,000 SCFW!
GAS FLOW RATE-SCFM
H4
FIGURE
-------
CITRATE PROCESS
TOTAL CAPITAL INVESTMENT COSTS
S02 REGENERATION SECTION
$10 Million
Costs: Mid 1974
10,000 Ibs/hr
SULFUR DIOXIDE RATE
OF REGENERATION SYSTEM (Ibs/hr)
145
FIGURE 9.4
-------
Table 9,2. CITRATE PROCESS
CAPITAL AND TOTAL ANNUAL COSTS
.e-
CTi
TOTAL CAPITAL INVESTMENT
Annual Cost
A. Direct Operating
1. Power
2. Water Process
3. Water Cooling
4. Citric Acid
5. Soda Ash
(Na2C03)
6 . Kerosene
7. Natural Gas
8. Alumina Catalyst
9. Labor
10. Maintenance
11. Taxes & Insurance
TOTAL ANNUAL DIRECT
OPERATING COST
B. Annual Capitalized
Cost
C. Net Total
Annualized Cost*
$
$/SCFM
$ /Annual ton
S02 Removed
$/yr/ton
S0~ Removed
z
$/yr/ton
SO2 Removed
70,000
4,640,000
66
169
96,300
25,700
27,600
82,500
13,800
99,000
249,300
3,500
210,200
188,300
116,000
1,109,500
41
610,200
1,038,800
38
Gas Flow Rate
100,000
5,630,000
56
143
137,600
35 , 300
39,500
117,800
19,600
141,500
356,200
4,900
227,800
230,400
140,800
1,446,200
37
740,300
1,312,100
33
- SCFM @ 1% SO
200,000
8,150,000
41
104
275,100
70,600
79,900
235,600
39 , 300
283,000
712,300
9,900
227,800
326,000
203,800
2,463,300
31
1,071,700
2,091,800
27
2
300,000
10,200,000
34
86
412,700
106,000
118,400
353,400
59,900
424,500
1,068,400
14,800
280,300
408,000
255,000
3,501,400
30
1,341,300
2,835,600
24
-------
&-3. CITRATE PROCESS
TOTAL SYSTEM CAPITAL AND ANNUAL OPERATING COSTS
A. TOTAL CAPITAL INVESTMENT
1. Gas Conditioning
2. Citrate Scrubbing &
Sulfur Recovery
TOTAL
B. TOTAL ANNUAL
OPERATING COSTS
1. Gas Conditioning
2. Citrate Scrubbing
& Sulfur Recovery
TOTAL
C. NET TOTAL
ANNUALIZED COSTS*
1. Gas Conditioning
2. Citrate Scrubbing
& Sulfur Recovery
TOTAL
$/ Annual ton
SO 2 Removed
$/yr/ton
SO 2 Removed
$/yr/ton
SO 2 Removed
Gas Flow Rate - SCFM @ 1% S02
70,000
1,800,000
4,640,000
6,440,000
234
273,100
1,109,500
1,382,600
50
321,100
1,038,800
1,359,900
49
100,000
2,275,000
5,630,000
7,905,000
201
370,600
1,446,200
1,816,800
46
419,100
1,312,100
1,731,200
44
200,000
3,550,000
8,150,000
11,700,000
149
613,500
2,463,300
3,076,800
39
672,200
2,091,800
2,764,000
35
300,000
4,650,000
10,200,000
14,850,000
126
843,300
3,501,400
4,344,700
37
740,000
2,835,600
3,575,600
30
*Based on Corporate Tax Rate of 48%.
-------
148
-------
10.0 AMMONIA PROCESS
10.1 PROCESS DESCRIPTION
Sulfur dioxide removal from gas streams by ammonia-based scrubbing
has been studied intermittently by various groups since the 1880's when
a British patent was first issued to Ramsey. Ammonia-based processes
are not amenable to throwaway operation because of the cost of ammonia
and the solubility and nitrogen value (with chemical oxygen demand) of
ammonium salts. The first such experimental processes were operated to
yield a product of ammonium sulfate for eventual fertilizer use. Com-
mercial and experimental modifications have evolved over the years with
the greatest developmental emphasis being placed on methods of regenerating
fche scrubbing liquor to reduce operating costs and produce various
useful products. Among the more widely studied methods of regenerating
2
the scrubbing liquor are:
1) thermal stripping to yield primarily sulfur dioxide,
2) oxidation to yield primarily ammonium sulfate,
3) disproportionation to yield ammonium sulfate and
elemental sulfur
4) acidulation to yield sulfur dioxide and an ammonium salt.
Recent U.S. work by the Tennessee Valley Authority and the Environ-
^ental Protection Agency has been directed toward a variation of the
Hixson-Miller scheme.3 The process employs acidulation of the scrubbing
liquor with ammonium bisulfate (ABS) to release sulfur dioxide, followed
by crystallization and decomposition of the resulting ammonium sulfate
to yield recyclable ABS and ammonia.4'3 This process appears to have
8everal advantages over other regenerable ammonia scrubbing schemes.
The energy requirements are lower than those for thermal stripping.
^lic acidulation and decomposition also avoid marketing problems
as8ociated with the large-scale production of ammonium sulfate or other
^desirable coproducts. The sulfur dioxide is a versatile product that
Can be liquefied and sold directly or converted into sulfuric acid or
®lemental sulfur.
' Of the many alternative regenerable ammonia scrubbing processes,
the ammonia absorption-ABS acidulation process was chosen as the most
for evaluation in this study. The overall process is summarized
149
-------
in the following paragraphs with implied reference to the process flow
diagram presented in Figure 10-1. The flow diagram essentially depicts
the process being piloted by TVA and EPA with appropriate modifications
to represent a complete commercial system.
1) Sulfur Dioxide Absorption — The S02-laden gas, which has been
cleaned of particulate material and cooled to about 125°F (52°C),
enters the absorber and flows countercurrently to an ammoniacal solution.
The absorber contains four or more valve trays, each of which is fed by
a scrubbing solution of different concentration. The liquor composition
is controlled to maximize absorption of SO,, while minimizing the loss of
NH« and the undesirable formation of an opaque fume in the cleaned gas.
The product liquor is withdrawn from the bottom stage for regeneration.
2) Acidulation and Stripping — The absorber product liquor is
pumped to a mixer and blended with ammonium bisulfate. The solution
then flows to a reactor where the product liquor and ABS react to form
ammonium sulfate and release SO^. Liquor from the acidulation reaction
is fed to a packed column where a countercurrent flow of air strips the
remaining S02 from the solution and blends it with offgas from the
reactor to form a concentrated SCL product stream.
3) Regeneration of ABS — The acidulated and stripped liquor is
pumped to an evaporator-crystallizer to concentrate and crystallize the
ammonium sulfate. The sulfate crystals flow to a belt filter and dryer
which reduce their moisture content to less than 1 percent. They are
then conveyed to a decomposer where ammonium sulfate decomposes at high
temperature to ammonia and ABS. The ammonia is returned to the absorption
system, and the ABS is recycled to the acidulation mixer.
4) Sulfate Purge -- A side stream is removed from the acidulated
and stripped liquor for purge requirements. The purge removes excess
sulfate formed in the absorber and controls other solid impurities in
the system. Limestone is added to the purge stream for neutralization.
The slurry is settled in a thickener, and the solids are removed as a
cake from a drum filter. The clear liquor is returned to the absorption
system.
150
-------
K.O. DRUM
ABSORBER
Ln
ACIDULATION
REACTOR--
PURGE
FILTER
PURGE
REGENERATOR
& THICKENER
MAKEUP
WATER
EVAPORATOR
CRYSTALLIZER
so2
TO ACID
PLANT
STRIPPING
AIR
SMELTER
GAS FROM
CONDITIONING
SECTION
MAKEUP
AMMONIA
STEAM-
DECOMPOSER
AMMONIA SCRUBBING PROCESS - ABS ACIDULATION
FIGURE 10-1
-------
10.2 PROCESS AND OPERATING CONSIDERATIONS
10.2.1 Absorption Section
The ammonia scrubbing process utilizes aqueous ammonia and ammonium
salt solutions to absorb sulfur dioxide from a gas stream. The net
absorption reactions are
HH4HS03 + NH3 * (NH4)2S03
30., + S00 + H00 * 2NH.HSO..
t .£ 3 22 43
Part of the sulfite is oxidized to sulfate primarily by the reaction
)0S00 + 1/2 00 *
The most significant operating parameters in the absorption step
have been shown to be:
1) solution temperature,
2) total concentration of S02 and NH3 in solution,
3) concentration of individual ammonium salts (sulfite,
bisulfite, and sulfate) which also determines pH,
4) ratio of liquid to gas flow,
5) type of internal column construction.
The important operating considerations of the abosrber are related
to the vapor-liquid equilibria of the system and the approach to equiU
brium conditions. In 1935 Johnstone obtained experimental data which
fit theoretically-based correlations for the vapor pressure of S02> NH
and water above ammoniacal solutions of varying composition, pH, and
temperature. Additional data and correlations were published in the
1950's and early 60's by Russian and English scientists.7'8'9'10
In 1974, Griffin recast the original Johnstone correlations and
showed that the equilibrium absorption of SO, is enhanced by
1) decreasing the solution temperature,
2) minimizing the total S02 concentration in solution, '
152
-------
3) minimizing S/Ca. the molar ratio of SO, in solution as ammonium
sulfite and bisulfite to ammonia as sulfite and bisulfite.
Ammonia losses are reduced by
1) decreasing the solution temperature,
2) minimizing the total NH3 concentration in solution,
3) maximizing S/C .
The vapor pressure of water in this system follows Raoult's law to
a good approximation. Although conflicting opinions have been reported,6'7'8'10
the sulfate concentration seems to have little effec? on SO and NH
vapor pressures in the range of sulfate concentration normally encountered
in scrubbers. Johnstone and Chertkov have also reported compatible
data which show that the solution pH is well correlated in the range of
4-8 to 6.6 with the bisulfite-sulfite ratio as the only parameter.
Obviously, the favorable conditions stated above are not all compatible
and must be optimized along with other system parameters.
The primary factors affecting the approach to equilibrium in the
absorption column are the internal absorber construction and the liquid
to gas flow ratio, L/G. Absorption of S02 in ammoniacal solutions has
een satisfactorily accomplished in both packed towers and tray columns.
11 the older units treating Cominco's zinc roaster and acid plant waste
gases at Trail, B.C.12*13 and Olin Mathieson's acid plant tail gas at
asadena, Texas, wood-slat packing was used in one or more stages with
'G's ranging from 16-32 gal/1000 CFM. Treating inlet gas concentrations
°.9 to 5.5 percent S0_, removal efficiencies of 92 to 97 percent were
Sieved. In a Cominco unit treating 0.75 percent S02 offgas from a
6Qd sintering plant, an L/G of 8 to 10 was used in the lower stages and
4 t- 12
c° 5 in the top stage to achieve an efficiency of 87 percent.
In the more recent TVA-EPA pilot plant work, multistage marble bed
Valve tray arrangements have been used. The best apparent combination
°r inlet gases of approximately 0.25 percent S02 is a four-stage valve
ay configuration, which allows flexibility with respect to the control
absorbent composition at different points in the column and achieves
6rall S02 removal efficiencies of about 90 percent.
In the process shown in Figure 1, makeup ammonia and ammonia
vcled from the regeneration unit are absorbed into solution and added
153
-------
to the second stage from the bottom. Ammonia stripped from the lower
two stages is captured in the water feed to the top stage and in the
weak salt solution overflowing to the third stage below. Product liquor
is withdrawn from the bottom of the column and a portion is recirculated
to the second stage for pH control. L/G's of about 10 at the bottom and
middle stages and 5 at the top stage have been used at the TVA-EPA pilot
4
plant.
Listed below are several general operating observations that follow
directly from the above discussion of controlling parameters:
1) Increasing the salt concentration of the bottom product
increases liquor concentration throughout the column and thus
increases SO,, and NH- vapor pressures above the top stage. A ^
water wash on the top stage can be used to control NH- losses.
2) A liquor pH below 5.6 allows essentially no S02,absorption;
a pH above 6.8 results in very high NH, losses. The liquor
pH on the top tray must be _^ 6.1 to control NH~ emissions to
50 ppm or less.
3) Temperatures must be kept relatively low to minimize NH-
losses and maintain a favorable equilibrium for S02 absorption.
Reported operating values range from 77°F to 125°F .
With respect to heat effects of a system scrubbing 100°F, 1 percent
S09 acid plant tail gas maintains a fairly constant temperature of 77°F
13
with no external cooling since the heat of reaction generated in the
system is balanced by humidification cooling of the dry gas. Heat
effects can be significant with higher S02 concentrations and a saturated
inlet gas stream, however. Depending on the ultimate product mix of
bisulfite, sulfite and sulfate, the net heat of reaction of the system
can vary from about 56,000 Btu/lb mole SO- absorbed to more than 210,000
Btu/lb mole S02 absorbed. A practical value occurring with a product
S/C of 0.8 and 13 percent oxidation of SO, to sulfate is about 89,000
3. £,
Btu/lb mole S02- After credit for gas humidifcation, if applicable,
this heat release may still require additional cooling between absorpti011
stages in the column to maintain the liquor temperature at an acceptable
value.
A serious problem which has been encountered in moat ammonia scrubb*
systems is the formation of an opaque fume in the exit gas stream. The
fume is partially attributed to gas-phase reactions of ammonia, S0» a°
154
-------
water forming ammonium sulfite, ' ' ' ' which are not prevented by a
mist eliminator. Carryover of salt through the mist eliminator may also
play an important role. The fume is objectionable as an opacity and
particulate emission problem and as a contributing factor to ammonia
loss. Cominco reported adequate control of the fume when operating with
liquor temperature of 77 °F and a clean inlet gas (low particulate
13
concentration and no acid mist) . Olin Mathieson operated with a pH
control system which reportedly eliminated the fume, but the exact pH
14
limitations are not reported. In the USSR a wet electrostatic pre-
cipitator is reportedly used at the top of ammonia scrubbers to control
the fume.16
At the TVA-EPA pilot plant, the fume has been controlled to 5
Percent opacity or less when operating with a liquor temperature of
about 120°F by using a prewash to remove particulates and SO-,, main-
taining a low salt concentration on the top stage, and reheating the
exit gas 10 to 20°F above the temperature required to dissipate the
steam plume. Even with these restrictions, the fume will often reform
outside the absorber on humid days.
10.2.2 Regeneration Section
For the complete regeneration process shown in Figure 10-1, the net
factions are
ACIDULATION
2S03 + 2NH4HS04
DECOMPOSITION
(NH4)2S04 & NH4HS04 + NH3 +
T^e net products of this regeneration system are water and ammonia,
wl*ich are recycled back to the absorption section, S02, and a small
PUrge of ammonium sulfate.
155
-------
The ABS acidulation scheme is based on a U.S. patent granted to
3
Hixson and Miller in 1946, but has not yet been applied commercially.
Acidulation of the absorber product liquor has been practiced by Cominco
since the 1930's by adding sulfuric acid to yield a net production of
12
S0~ and ammonium sulfate. Acidulation with nitric and phosphoric
acid, also leading to the production of S09 and fertilizer, has been
2
reported in Czechoslavakia and Romania, respectively.
The uncertainty of a large market for ammonium-based fertilizers
limits the applicability of these otherwise proven processes in the
United States so the TVA-EPA pilot program was designed to concentrate
primarily on the ABS acidulation scheme. Since the regeneration system
has not been fully integrated, sulfuric acid has been used in the pilot-
scale tests completed to date. The acid is metered to the acidulator to
give an acid ion to ammonia ratio of 1.2 to 1.5. The retention time of
the acidulator can be varied from 0 to 30 minutes. The solution leaves
the acidulator and flows countercurrently to stripping gas (air at about
3
30 ft /gal solution) to remove the chemically released SO,,. Under these
conditions, up to 97 percent release of the SO,, in the bisulfite-sulfite
liquor has been accomplished in the combined acidulation-stripping
4 5
operation. ' The product gas contains 20 to 30 percent S02, varying
according to the concentration of the absorber product liquor and the
gas to liquor ratio in the stripper.
Laboratory tests have shown that ammonium bisulfate acts very
similarly to sulfuric acid in the acidulation step and should give
comparable results. The chemical equations stated above show that a
high bisulfite to sulfite ratio (high S/C ) in the absorber product
3.
stream reduces the amount of ABS required for acidulation and thus
reduces the equipment size of the entire regeneration cycle. This
factor must be weighed against S/C restrictions in the absorber.
a
The product liquor from the stripper is pumped to an evaporator-
crystallizer where water is evaporated and ammonium sulfate crystals
formed. Provisions must be made to remove the S07 not stripped from
liquor in the stripper. This residual S02 is released with steam in
evaporator. When the steam is condensed, part of the S0? is dissolved
in the water and recycled to the absorption aection, and the excess
be bled from the condenser to the product S02 stream or elsewhere.
156
-------
The two-phase mixture from the crystallizer is sent to a belt
filter where the crystals are dewatered to about 5 percent moisture.
A gas-fired dryer lowers the water content further to about 1 percent,
Yielding free-flowing crystals, approximately 70 percent of which are
retained on a 35-mesh screen.
The total water that must be removed from the crystals increases
with a decrease in the salt concentration leaving the absorber. For
example, with a product liquor concentration of 12 moles NH3 (as sulfite
and bisulfite) per 100 moles total water
-------
An electric furnace decomposer is to be installed at the TVA-EPA
pilot plant in late 1975 for the continuation of that project. When
this decomposer and the associated ABS acidulation process are fully
integrated into the pilot plant, important operating variables of the
regeneration cycle can be better defined.
10.3 DEVELOPMENT STATUS
The ammonia scrubbing-ABS acidulation process, which is being
jointly developed by TVA and EPA and is shown schematically in Figure 10-1>
3
was initially patented by Hixson and Miller in 1946. Although the
process has never been operated in an integrated fashion on a signifi-
cant scale, the individual unit operations have all been proven to be
technically feasible in large industrial applications.
Scrubbing SO^ from waste gases with ammoniacal solutions has been
practiced commercially by Cominco in Trail, B.C., by Olin Mathieson in
Texas (using the Cominco system), and in Romania, Russia, Czechoslavakia»
20 21
Germany, Japan, and France. ' Acidulation of the absorber liquor to
release SO™ and form an ammonium salt has been practiced commercially in
20 21
Germany using t^SO, and in Romania using H_PO,. ' Acidulation with
ammonium bisulfate has not yet been practiced on a large scale. De-
composition of ammonium sulfate to yield ABS and ammonia was accomplished
successfully on a large scale in a U.S. Government sponsored program
during World War II. More recently in France, a direct-fired decompose*
has been developed to produce ABS from ammonium sulfate for metallurgies1
21
uses. The success of this unit is not yet known.
The Tennessee Valley Authority and the Environmental Protection
Agency began studies on an ammonia scrubbing program at TVA's Colbert
Power Plant in 1968. The pilot plant work accomplished to date has been
concentrated primarily on the absorption process. Extensive tests have
been run on a 3000 CFM unit to establish important operating parameters ^
and develop solutions to process difficulties such as the fume emission-
Preliminary regeneration tests have also been made.5 Complete integrati0
of the ABS acidulation and ammonium sulfate decomposition steps into
pilot plant should be accomplished in late 1975 or 1976.
158
-------
Although the individual unit operations have been applied in
industrial systems, it is obvious from the preceding sections that several
features have not yet been satisfactorily demonstrated (e.g., consistant
control of the fume emission from the absorber). In an integrated system
additional problems are likely to be encountered. Despite the extensive
work accomplished to date, the ammonia scrubbing-ABS acidulation process
cannot be considered fully developed for commercial application until
extended operational tests of an integrated system are completed at the
TVA-EPA pilot plant.
10.A DESULFURIZATION EFFICIENCY
A high desulfurization efficiency of weak or strong S02 streams can
be achieved with the ammonia absorption system provided adequate absorption
capacity is provided and the salt concentration and solution temperature
are properly controlled, particularly at the top stage. The older
Cominco units exhibited high efficiencies in various applications as shown
in the table below.
Type of Gas
Lead Sintering
Plant12
Zinc Absorption
Plant12
Acid Plant12'13
Gas Volume
ft3/min
150,000
20,000
50-60,000
Inlet S02
%
0.75
5.5
0.9-1.0
Outlet S02
%
0.10
10.2
0.07-0.08
S02 Removal
Efficiency
%
85
95
92
The TVA-EPA pilot plant has generally been operated with an S02 removal
e*ficiency of 90 percent or higher, with a feed concentration of 0.2 to
'0.3 percent S02 and an exit gas concentration of 200 to 300 ppm SCy
159
-------
From these experiences, it is evident that desulfurization efficiencies
of 90 to 95 percent are technically achievable. Since the salt concentration
and solution temperature also affect ammonia losses from the absorber,
as well as energy requirements in the regeneration section, the costs
associated with these variables are essential considerations in determining
the efficiency that can be achieved practically in a particular application-
10.5 SULFITE OXIDATION
The cause and effect relationships of sulfite oxidation in the
ammonia scrubbing process are not well established although extensive
studies have been conducted, primarily in Russia. At least three separate
mechanisms can contribute to oxidation as represented by the following
reactions:
3S02
1/2
The first reaction is insignificant at low temperatures unless catalysts
I I C^.
are present in the system. Several metallic salts such as Mn , V ,
l I
and Fe , which are commonly found in fly ash, may accelerate this
mechanism. A prewash before the absorber should minimize the concentrat*0
of these catalysts and should also inhibit the second reaction by sc
S03 from the gas. The third reaction is probably the most significant
and is directly related to the oxygen content of the flue gases and the
salt concentration of the scrubbing liquor.
In the Cominco scrubbers, where ammonia sulfate is considered a
product, oxidation is relatively unimportant. Sketchy information from
European ammonia scrubbing processes indicates oxidation levels of 10 c
30 percent of the inlet S02. In the TVA-EPA pilot plant approximately
10 to 15 percent of the S02 is typically oxidized to sulfate with an
average of 13 percent. Scattered data from this pilot plant appear to
indicate that increasing the S/C increases oxidation. This is in
cl
qualtitative agreement with Chertkov's correlation of industrial data
160
-------
which indicates that oxidation rate is proportional to (S/C ) to the
19 3
sixth power. On the other hand, the chemical equation shown above
indicates that oxidation of ammonium sulfite is the appropriate mechanism,
and thus for a constant S09 concentration, decreasing S/C would encourage
. "•
oxidation. Operating experience reported by Olin Mathieson confirms
that low pH (high S/C ) inhibits rather than promotes sulfate formation.
3.
The large number of variables and possible mechanisms acting on the
system obviously contribute to this conflict of opinion. Additional
study is needed in this area.
Newall published data which indicate that the presence of ammonium,
sulfate in the scrubbing solution significantly increases the vapor
Pressure of S02 and NH3 and thus adversely affects the scrubbing efficiency
and ammonia conservation. Newall's data correspond to a scrubber operating
with low salt concentration and over 50 percent oxidation. However,
Chertkov7 also indicated an adverse effect of sulfate concentration on
S02 vapor pressure, but his data emphasized dilute solutions and solutions
containing higher than normal ratios of bisulfite to sulfite. There are
apparently no data reported that indicate a significant sulfate effect
on absorption equilibria with the solution conditions encountered in a
Practical scrubber installation, but additional study in this area would
also be prudent.
With the ABS regeneration system, oxidation definitely increases
the purge requirements. To maintain a material balance in the regeneration
Action, sulfate must be purged to the extent that it is formed in the
absorber. Figure 10-1 shows this purge accomplished by removing a
sidestream from the stripper effluent and reacting with limestone to
yleld a gypsum cake for disposal and a clear liquor for recycle to the
aWrber. If a small market for ammonium sulfate exists, a feasible
alternative is to purge ammonium sulfate crystals prior to decomposition.
10-6 GAS PRETREATMENT
Waste gas from any process must be cooled to a reasonable temperature
Pt*or to being fed to an absorption tower to avoid excessive evaporation
°f the scrubbing solution. With ammonia scrubbing the "reasonable
tetnPerature" is established as a trade-off between the absorption capacity
an
-------
is exothermic, some humidification cooling of a hot, dry gas stream can
be allowed in the tower to eliminate the need for cooling of the liquor
between stages. For example, operating experience at Cominco's acid
plant tail gas treatment system showed that humidification cooling of
the 104°F, dry inlet gas stream resulted in 77°F tower operation with no
13
external cooling. Liquor temperatures of up to 130°F (after pre-
cooling of the gas) have been utilized at the TVA-EPA pilot plant with
acceptable SCL absorption and NH3 loss. The gas cooling and conditioning
system for regenerable systems which is described in Appendix A would
provide acceptable pre-cooling for the ammonia scrubbing process, but
interstage cooling of the liquor would probably be necessary with stronger
SO, streams (perhaps 1 percent and higher) since the inlet gas would be
saturated with water.
Pretreatment would also be necessary in most applications to minimize
SO and particulate concentrations in the tower. Sulfur trioxide as
well as certain metallic oxides in the flue gas would promote oxidation
in the tower. Particulates such as fly ash can also cause problems in
the regeneration section by adversely affecting the production of cry-
stalline ammonium sulfate and by catalyzing the decomposition of ammonia
4
during the thermal decomposition of ammonium sulfate.
In the TVA-EPA pilot plant, a low-pressure-drop venturi scrubber is
used to humidify, cool, and remove solids from the inlet gas stream.
The inlet gas is a slipstream from a coal-fired boiler and enters at
about 295°F with a dust loading of 4 to 6 grains per dry standard cubic
foot and an S02 concentration of 0.2 to 0.3 percent. The venturi is
operated with a pressure drop of 5 to 6 inches of water and a liquid to
3
gas ratio of about 10 gal/1000 ft . Under these conditions approxi-
mately 90 percent of the fly ash and 10 to 20 percent of the SCK are
removed. Trace amounts of chloride (35 ppm average) in the inlet gas
are also completely removed. The gas is cooled to saturation humidity
at 120 to 130°F . Satisfactory absorption tower operation is achieved
with this degree of pretreatment. While the experience is not directly
relatable to smelter offgases, it exemplifies the relative degree of
pretreatment considered necessary for the ammonia absorption system.
162
-------
The conditioning system described for regenerable processes in
Appendix A should therefore provide adequate cooling and particulate
removal. The important additional cost factor for the ammonia system is
the probable requirement for interstage cooling to maintain the tower
temperature at 130°F or lower.
10.7 PARTICULATE REMOVAL CAPABILITY
Pretreatment requirements with respect to particulate removal were
discussed in the preceding section. There is no detailed information
available on the operation of the absorber without prior removal of
Particulates although the Russians are reported to favor this design to
avoid the expense of a separate dust scrubber or precipitator.21 The
ash is said to be easily filterable from the absorber liquor.
At the TVA-EPA pilot plant a yellow solid, tentatively identified
Petrographically as a homogeneous iron-ammonia-sulfur compound is
flormally present in the absorber liquors even though 90 percent or
greater particulate removal is accomplished prior to the absorber. It
w°uld appear that precipitation of solids in the absorber could be a
Problem especially if no pretreatment were provided. The solids tend to
Settle in the product storage tank. In contrast to the Russian claims,
attempts to remove the solids by filtration have been unsuccessful at
"e pilot plant because the precipitated solids and fly ash form a
Gelatinous, thixotropic material that blinds the filter media. In a
Belter application the solids may not behave in this manner.
The extraneous solids that do not settle in storage and process
essels would normally be controlled by the purge stream. Depending on
e dust load of the inlet gas and the allowable concentration in the
Drubbing liquor, the purge requirement to control solids may not be
nsistent with the normal purge rate required to maintain a sulfate
a-Unce. The solids may also interfere with ammonium sulfate crystalli-
ation and could feasibly catalyze the decomposition of ammonia in the
ern»al decomposer. Without sufficient operating experience to access
6Se potential problems, the particulate removal capability of the
°cess remains in question.
163
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10.8 PROCESS ENERGY REQUIREMENTS
Power requirements in the conditioning and absorption sections of
the ammonia process should be comparable to other S02 scrubbing processes
since the standardized conditioning system described in Appendix A is
applicable to the ammonia system, and flue gas pressure drop through the
valve-tray absorber is moderate. Pumping power requirements will be
slightly increased with stronger SCL streams which require interstage
cooling of the absorbent. The estimated total power requirement for the
absorption section (not including the separate gas-conditioning section)
is about 4kW/1000 SCFM.
In the regeneration section, the predominant energy requirements
are steam for the evaporator and decomposer and electricity for the
decomposer. With a typical absorber product liquor C value of 12, the
a 4
evaporation requirement is about 2.5 to 2.9 Ib water per Ib S0? product.
Evaporation requirements increase with a more dilute product liquor.
18
Decomposer design studies based on the Plancor 1865 project show
requirements of approximately 350 kWh electric power and 400 Ib super-
heated steam per ton of ammonium sulfate fed to the decomposer (based on
approximately 85 percent decomposition of ammonium sulfate to ABS).
With a typical absorber product liquor S/C of 0.8, these figures convert
3.
to 0.45 kWh electricity and 0.5 Ib superheated steam per Ib of S02
product.
10.9 RETROFITTING REQUIREMENTS
Space requirements for the ammonia process should not be signifi-
cantly different from the other S0? scrubbing processes. The gas con-
ditioning and absorption sections will consist primarily of vertical
2
vessels, pumps, and piping, requiring an estimated 8-12,000 ft for a
system capacity on the order of 100,000 SCFM. To tie into existing
ductwork and keep fan power requirements to a minimum, these sections
would be installed as close as possible to the existing plant and stack.
The largest space requirement in the regeneration section will be
1 8
the decomposer. In the Plancor operation, four furnaces, each 16 ft
in diameter by 10 ft deep, were designed for a total feed rate of about
550 tons (NH4)2S04 per day. By contrast a 100,000 SCFM absorption
system treating 2 percent S02 gas will require a decomposer capacity of
approximately 600 tons (NH^)SO, per day. The total space requirement of
a regeneration system of this capacity, including the purge system shown
164
-------
in Figure 10-1, is estimated to be 12-15,000 ft2. This requirement is
obviously related to both the volume of gas treated in the absorption
system and the sulfur dioxide concentration. It is not imperative that
the regeneration facility (especially the purge system) be adjacent to
the existing plant, but it would normally be sited as close as possible
to minimize pumping costs.
10.10 PROCESS COSTS
Published cost data on the ammonia system are limited to estimates
which have been based primarily on the TVA-EPA pilot plant and on the
literature studies conducted in preparation for the TVA-EPA project. In
1970, TVA published an extensive literature survey and study of the
ammonia scrubbing process including cost estimates for application to
Power plant stack gases yielding ammonium sulfate for fertilizer pro-
0*3
duction. M. W. Kellogg utilized these estimates as a basis for another
f\ i
study in 1971. More recently the Lummus Company prepared for the
Control Research Association, Inc., an engineering estimate for
aPplication of the ammonia scrubbing-ABS acidulation process to 1 percent
s°2 copper reverberatory furnace gases. These estimates, supplemented
by standard chemical engineering cost estimation techniques, ' provided
"e basis for developing capital costs for the absorption and regeneration
Actions.
Capital costs versus gas flow for the absorption section are pre-
Sented in Figure 10-2. Ammonia scrubbing will in general require a
arger capital investment than the "standardized" absorption systems
Presented in Appendix B because of the interstage liquor cooling requirements.
only one curve is shown in Figure 10-4, the capital cost of the
section will actually be somewhat dependent on S09 concentration
j
11 the gas stream. Capital cost of the regeneration and purge section is
Presented in Figure 10-5 on the basis of S02 handled in Ibs/hr. The
Ccuracy of both of these correlations is of course limited by the lack
e*perience with a large scale integrated ammonia-ABS system.
The regeneration and purge costs for various SO* rates have been
with the capital cost of the S00 absorption section to provide
ara
al
th
6 Parametric capital cost relationships presented in Figure 10-2.
annual operating costs have been developed from estimates of the
fr>
C8y, chemical make-up, and operating labor requirements of the system,
Ined with maintenance, taxes and insurance expressed as a percentage
165
-------
of invested capital. These parametric curves are presented in Figure
10-3. Unit bases and costs are shown in Table 10-1.
For comparative purposes, Table 10-2 provides a capital and operating
cost breakdown for this process applied to a range of gas flows containing
1 percent S0~. Table 10-3 estimates the total system cost including gas
cooling and conditioning as described in Appendix A and an auxiliary
sulfuric acid plant as described in Appendix C. This table allows
direct comparison with similar system costs for the other S02 control
processes under evaluation.
10.11 ADVANTAGES AND DISADVANTAGES
Advantages of the ammonia absorption-ABS acidulation process are as
follows:
1) The process can achieve high efficiencies of SO,, removal over
a wide range of S02 concentration;
2) Reaction products in the absorber are soluble salts which
avoid the scaling and plugging problems of some alternative
processes;
3) The integrated absorption-regeneration system produces a
concentrated S02 stream which can be used to produce sulfuric
acid, elemental sulfur, or liquid SO,,;
4) A potential for sale of small amounts of ammonium sulfate may
eliminate the need for purge treatment equipment and minimize
the penalty of sulfite oxidation;
5) The ABS acidulation scheme requires only one-half to two-
thirds as much energy as the Johnstone steam-stripping process
and avoids the disproportionation reactions and sulfite purge
encountered in other regeneration processes following ammonia
absorption of SO™.
Disadvantages of the process are:
1) Ammonia volatility may limit the minimum level of SO™ emission
to 200-300 ppm for practical operation;
2) Ammonia may also be destroyed during the decomposition of
ammonium sulfate;
3) Satisfactory elimination of the fume emitted from the absorber
has not yet been completely achieved in the latest EPA-TVA
work;
4) The ABS regeneration scheme has not yet been demonstrated in
an integrated system. Until this is achieved at least on a
pilot scale, the process will probably not be commercially
viable.
166
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10.12 REFERENCES
1. Ramsey, "Use of the NH3-S02-H20 System as a Cyclic Recovery Method,"
British Patent 1,427 (1883).
2. Slack, A.V., "Sulfur Sioxide Removal from Waste Gases," Park Ridge
New Jersey, Noyes Data Corporation, 1971 (Pollution Control Review '
No.4).
3. Hixson, A.W. and R. Miller, "Recovery of Acidic Gases," U.S Patent
2,405,747 (August 13, 1946).
4. Tennessee Valley Authority, "Pilot-Plant Study of An Ammonia
Absorption-Ammonium Bisulfate Regeneration Process, Topical
Report Phases I and II," U.S. Environmental Protection Agency,
Environmental Protection Technology Series, EPA-650/2-74-049-a
(June 1974).
5. Holliden, G.A., and C.E. Breed, "TVA-EPA Pilot-Plant Study of the
Ammonia Absorption-Ammonium Bisulfate Regeneration Process,"
paper presented at the Flue Gas Desulfurization Symposium, Atlanta,
Ga., Nov. 4-7, 1974 (proceedings to be published).
6. Johnstone, H.G., "Recovery of Sulfur Dioxide from Waste Gases:
Equilibrium Partial Vapor Pressure over Solutions of the Ammonia-
Sulfur Dioxide-Water System." Ind. Eng. Chem., 27(5), May 1935,
pp 587-593.
7. Chertkov, B.A. and N.S. Dobromyslova, "The Influence of Traces of
Sulfate on the Partial Pressure of S02 over Ammonium Sulfite-
Bisulfite Solutions." J. Appl. Chem. USSR, 37(8), 1964, pp 1707-1711.
8. Chertkov, B.A., D.L. Puklina, and T. I. Pekareva, "The pH Values of
Ammonium Sulfite-Bisulfite Solutions." J. Appl. Chem. USSR, 32(6),
1959, pp 1417-1419.
9. Berdyanskaya, R.A., S.M. Golyand, and B.A. Chertkov, "On the Partial
Pressure of S02 Over Ammonium Sulfite-Bisulfite Solutions," J. Appl.
Chem. USSR, 32, 1959, pp 1978-1984.
l°. Newall, H.E., "Ammonia Process for Removal of Sulfur Dioxide from
Flue Gas," in Problems and Control of Air Pollution, F.S. Mallette,
ed., New York, Reinhold, 1955, pp 170-190.
11• Griffin, L.I., "Evaluation of Equations for Designing Ammoniacal
Scrubbers to Remove Sulfur Oxides from Waste Gases," U.S.
Environmental Protection Agency, Environmental Protection Technology
Series, EPA-650/2-74-035 (January 1974).
l2- King, R.A., "Economic Utilization of Sulfur Dioxide from Metallurgical
Gases," Ind. Eng. Chem., 42(11),(November 1950), pp. 2241-8.
l3- Burgess, W.D., "SO, Recovery Process as Applied to Acid Plant Tail
Gas," Chemistry in Canada, June 1956, pp 116-119.
167
-------
14. Lehle, W.W., "Processing of Waste Gases from Sulfuric Acid Plants,"
in The Manufacture of Sulfuric Acid, W.W. Duecker and J.R. West, ed.,
New York, Reinhold, 1959, pp 346-358.
15. Hein, L.B., A.B. Phillips, and R.D. Young, "Recovery of Sulfur
Dioxide from Coal Combustion Stack Gases," in Problems and Control
of Air Pollution, F.S. Mallette, ed., New York, Reinhold, 1955,
pp 155-169.
16. "Ammonia-Cyclic Scrubbing of SC^ from Flue Gas of Power Stations,"
Ministry of Chemical and Petroleum Machine Building, the State
Scientific Research Institute for Industrial and Sanitary Gas
Cleaning, Moscow, 1973.
17. Newcombe, G.M., "Technical Status Report, Ammonia Scrubbing Program,"
intralaboratory report of Control Systems Laboratory, U.S. Environ-
mental Protection Agency, April 6, 1972.
18. Engineer - Contractor's Report on Alumina-from-Clay Experimental
Plant (Plancor 1865) at Salem, Oregon, 1946.
19. Chertkov, B.A., "General Equations for the Oxidation Rate of Sulfite-
Bisulfite Solutions in the Extractioi
Chem. USSR 34(4), 1961, pp. 743-747.
Bisulfite Solutions in the Extraction of S02 from Gases," J.Appl.
20. Newcombe, G.M., M.A. Maxwell, and G.T. Rochelle (EPA Control Systems
Laboratory), "Control of Sulfur Dioxide Emission from Stack Gases:
Discussion of Promising Soluble-Base Aqueous-Scrubbing Processes,"
unpublished paper presented at Delaware Valley Section Meeting, A.I.Ch-
Philadelphia, Pa., March 21, 1972.
21. Griffin, L.I., personal communication, August 1974.
22. Holliden, G.A., N.D. Moore, P.C. Williamson, and D.A. Denny, "Removal
of Sulfur Dioxide from Stack Gases by Scrubbing with Ammoniacal
Solutions: Pilot-Scale Studies at TVA," Proceedings: Flue Gas
Desulfurization Symposium (1973), U.S. Environmental Protection Agency**
Environmental Protection Technology Series, EPA-650/2-73-038 (December
1973) pp 961-996.
23. Tennessee Valley Authority, "Sulfur Oxide Removal from Power Plant
Stack Gases", prepared for National Air Pollution Control Administrat
under Contract No. TV-29233A, 1970.
24. M.W. Kellogg Company, "Evaluation of S02~Control Processes, Task ®°
Final Report". U.S. Environmental Protection Agency (APTD-807), Oct°
1971.
25. Smelter Control Research Association,,Inc.,"Report to U.S. Bureau
of Mines on Engineering Evaluation of Possible High Efficiency
Soluble Scrubbing Systems for the Removal of S02 from Copper
Reverberatory Furnace Flue Gases." B.O.M. Contract No. S0133044,
March 1974.
168
-------
26. Peters, M.S. and K.D. Timmerhaus, Plant Design and Economics for
Chemical Engineers, 2nd edition McGraw Hill Book Company, New
York, 1968.
27. Guthrie, K.M., Process Plant Estimating Evaluation and Control,
Craftsman Book Company of America, Solvanca Beach, Calif. 1974.
169
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Table 10.1 AMMONIA PROCESS UNIT USAGE AND COST DATA
A. Chemical Utilities
Basis
Unit Cost
Power a) S02 Absorption
b) S00 Regeneration
Water a) Process
b) Cooling
Ammonia (anhydrous)
a) Loss from scrubber
b) Loss from purge
Limestone
Natural Gas
Steam a) 15 psig saturated
b) 15 psig, 700°F
4 kW/1000 SCFM
0.392 kWh/lbS02
absorbed
0.06 gal/lbSO.
absorbed
12.8 gal/lbS02
absorbed
50 ppm
0.0126 lb/lbS02
absorbed
0.25 lb/S02
absorbed
0.145 CF/lbSO,
absoroed
2.27 lb/S02
absorbed
0.45 lb/lbS02
absorbed
$0.015/kWh
$0.30/Mgal
$0.10/Mgal
$190/ton NH3
$ 4/ton limestone
$1.25/MCF
$0.80/M Ib steam
$1.25/M Ib steam
B. Operating Labor &
Maintenance
Labor: Absorption
S0_ Regeneration
Maintenance:
Taxes and Insurance
% man/shift
<75,000 SCFM
3/4 man/shift
>75,000 SCFM
2% man/shift
<10,000 Ib/yr
3% man/shift
>10,000 Ib/hr
4.0% TCI/year
.IL5LIS/ZSSL
$8/hr
C. Fixed Charges
13.15% TCI/year
Based on Capital Recovery Factor using 10% interest over 15 year
170
-------
AMMONIA SCRUBBING - ABS ACIDULATION
TOTAL CAPITAL INVESTMENT COSTS
S02 REMOVAL EFFICIENCY = 95%
.O MILLION
NOTE:
GAS COOLING AND CONDITIONING
NOT INCLUDED
100,000 SCFM
GAS FLOW RATE-SCFM
171
4.0% SO.
Costs: Mid 1974
FIGURE 10-2
-------
AMMONIA SCRUBBING - ABS ACIDULATION
TOTAL ANNUAL DIRECT OPERATING COSTS
S02 REMOVAL EFFICIENCY =
95%
$1.0 Million
NOTE: GAS COOLING AND CONDITIONING
NOT INCLUDED
100,000 SCFM
GAS FLOW RATE-SCFM
172
-------
AMMONIA SCRUBBING - ABS ACIDULATION
TOTAL CAPITAL INVESTMENT COSTS
S02 REGENERATION SECTION
S02 REMOVAL EFFICIENCY = 95%
$10 Million
Costs: Mid 1974
10,000 LB/HR
SULFUR DIOXIDE RATE,
IN REGENERATION SECTION (Ibs/hr)
173
FIGURE 10-4
-------
AMMONIA SCRUBBING - ABS ACIDULATION
(TOTAL CAPITAL INVESTMENT COSTS)
4-STAGE AMMONIACAL SCRUBBER
WITH LIQUOR INTERCOOLING
to
tO
O
O
LU
to
LU
1.0 MILLION
a.
<
O
O
Costs:
100,000 SCFM
GAS FLOW RATE-SCFM
174
FIGURE
-------
J.O.2- AMMONIA PROCESS
CAPITAL AND TOTAL ANNUAL COSTS
TOTAL CAPITAL INVESTMENT
Annual Cost
A. Direct Operating
1 . Power
2. Process Water
3. Cooling Water
4. Ammonia
5. Limestone
6. Natural Gas
7. Steam '
8. Labor
9 . Maintenance
10. Taxes & Insurance
TOTAL ANNUAL DIRECT
OPERATING COST
B. Annual Capitalized
Cost
C. Net Total
Annualized Cost*
$
$/SCFM
$ /Annual ton
S02 Removed
$/yr/ton
S02 Removed
$/yr/ton
SO 2 Removed
Gas Flow Rate - SCFM @1% SO,
70,000
5,900,000
84
214
376,000
1,000"
74,000
77,000
29,500
10,500
137,800
210,200
236,000
147,500
1,299,500
49
775,900
1,262,800
46
100,000
7,450,000
75
190
537,100
1,500
105,200
110,000
42,100
15,000
196,900
227,800
298,000
186,300
1,720,400
44
979,700
1,635,900
42
200,000
10,450,000
52
133
1,074,100
3,000
211,300
220,100
84,200
30,000
393,900
227,800
418,000
261,300
-
2,923,700
37
1,374,200
2,560,100
33
Based on Corporate Tax Rate of 48%.
300,000
13,000,000
AQ
HJ
110
1,611,200
4,500
317,000
330,200
126,300
45,100
590,800
280,300
520,000
325,000
=
4,150,400
35
1,709,500
3,451,700
30
-------
Table 10.3 AMMONIA PROCESS
TOTAL SYSTEM CAPITAL AND ANNUAL COSTS
A. TOTAL CAPITAL INVESTMENT
1. Gas Conditioning
2. Ammonia Scrubbing-
ABS Acidulation
3. Auxiliary Sulfuric
Acid Plant
TOTAL
B. TOTAL ANNUAL
OPERATING COSTS
1. Gas Conditioning
2. Ammonia Scrubbing-
ABS Acidulation
3. Auxiliary Sulfuric
Acid Plant
TOTAL
C. NET TOTAL.
ANNUALIZED COSTS*
1. Gas Conditioning
2. Ammonia Scrubbing-
ABS Acidulation
3. Auxiliary Sulfuric
Acid Plant
TOTAL
$/ Annual ton
SO- Removed
$/yr/ton
S02 Removed
$/yr/ton
SO2 Removed
70,000
1,800,000
5,900,000
1,450,000
9,150,000
333
273,100
1,299,500
199,500
1,772,100
66
321,100
1,262,800
248,000
1,831,900
67
Gas Flow Ral
100,000
2,275,000
7,450,000
1,820,000
11,545,000
293
370,600
1,720,400
234,100
2,325,108
61
419,100
1,635,900
302,800
—_ ^--I ?_: -— =.
2,357,800
60
:e - SCFM @1% S(
200,000
3,550,000
10,450,000
2,775,000
16,775,000
214
613,500
2,923,700
271,200
3,908,400
51
672,200
2,560,100
469,000
3,701,100
47
>2
300,000
4,650,000
13,000,000
3,600,000
21,250,000
180
843,300
4,150,400
475,000
5,468,700
46
740,000
3,451,700
605,200
4,796,200
41
-J
o\
-------
11.0 APPLICATION OF ABSORPTION-BASED SO, CONTROL SYSTEMS TO
WEAK S00 COPPER SMELTER REVERBERATORY FURNACES
11.1 GENERAL
Development of S02 absorption-based control processes has been
generally oriented towards power utility plant application. The flue
gases from such plants, whether coal- or oil-fired, are characterized by
1) large uniform gas volumes, 2) moderate temperatures of around 300°F,
3) low levels of S02—usually less than 0.3 percent, and 4) moderately
low levels of oxygen—5 percent or less.
Some limited pilot plant development work on nonferrous smelter gas
slip streams or smelter-simulated gas streams has been done, or is in
Progress, on the limestone, double alkali (both sodium and ammonia
alternatives), and citrate processes. Results have ranged from dis-
couraging for the limestone process, to promising for the citrate process.
Much additional work is still required.
Only three S09 absorption processes have had the benefit of full-
8cale commercial application in a smelter environment—the ammonia
Process usually producing ammonium sulfate as the end product, ASARCO's
dimethylaniline process, and the magnesium oxide process on a rever-
beratory furnace gas in Japan. ASARCO's DMA process is presently
Derating on copper converter gases at the ASARCO Tacoraa smelter. A
u*it has also been installed by Phelps-Dodge at the Ajo smelter but has
been plagued with mechanical problems and has not yet demonstrated its
CaPability on a weak SO™ gas stream.
It is only after many years of development and a number of demon-
stration installations in the utility area that the lime/limestone
8c*ubbing process has been able to demonstrate the reliability and
Process capability which a control process in an industrial environment
have. Investment and operating costs have continued to climb as
needs were established and equipment requirements were better
8Pecified. While much has been gained from this development period, it
^t be appreciated that it has taken place under the essentially uniform
c°n
-------
In considering the application of absorption-based SO^ control
systems to the weak S0« streams of nonferrous smelters, it is necessary
to give some thought to the differences which exist between smelter
operation and utility practice and the effect these differences may have
on the control process itself.
11.2 SMELTER REVERBERATORY FURNACE AND ROASTER OPERATIONS
In reverberatory furnaces, smelting of the copper-carrying ore
concentrate or calcine is accomplished by the combustion of large volumes
of fuel and air. The offgases are accordingly large in volume with low
concentrations of the sulfur oxides. The operating conditions are
established to produce from a given concentrate or calcine, a copper
matte with a composition which facilitates conversion to blister copper
during the converter step. However, there are a number of major factors
which directly affect reverberatory furnace operation and hence the
volume of the effluent gases and the concentration of S0~ in these
gases:
1) The mineralogy of the ore concentrates— Copper-bearing
ores, even after concentration, differ widely in their
mineral makeup. The levels of pyritic sulfur, silicate
makeup and content, etc., exercise a strong influence
on the readiness with which an ore will smelt and hence
the firing rate or fuel/charge rates necessary to achieve
the desired reactions. It is apparent, therefore, that
reverberatory operations may differ markedly between
different smelters processing different source ore
material. Some smelters process ore on a custom basis
and, as purchasing contracts are competitively negotiated
and ore sources change, operating conditions in their
reverberatories could change dramatically.
2) The type of reverberatory charge — Four of the domestic
copper smelters under review presently use multi-hearth
roasters and two use fluo-solids roasters to produce a
calcine charge to their reverberatory furnaces. Roasting
tends to reduce the sulfur eliminated in the reverberatory
and results in weaker S0? streams.
3) The method of charging the reverberatory furnace — A
reverberatory furnace is charged approximately every
15-20 minutes. Charging through the roof (side-wall
charging) is invariably associated with peaks in
both gas flow and S02 concentration. Charging by
Wagstaff guns, which disperse the concentrate or
calcine over the molten bath, tends to reduce emissions.
178
-------
4) Treatment of converter slag — It is common practice to
return converter slag to the reverberatory furnace and
this approach will tend to increase the sulfur oxide
emissions per unit of charge. In one smelter at least,
this practice is no longer followed and the converter
slag is treated in separate facilities.
5) Air infiltration — Air infiltration into the gas ductwork
system of a reverberatory furnace is a universal occurrence,
although the condition of the ductwork and equipment and the
operating practices exercise a significant influence on the
total amount of infiltration. However, additional dilution
of the reverberatory offgases is also related to the
necessity of periodically "blowing" the waste heat boiler
of deposited particulates carried over from the furnace.
Air is usually used, but in some smelters, the system makes
use of steam.
Theoretically, the offgases from multi-hearth roasters should be
relatively uniform with an S02 concentration of 5-10 percent, but operating
roasters are usually old and considerable air ingress is common. It is
also common practice to use air dilution to effect cooling of the effluent
Bases. Actual operating conditions of the roasters and sulfur elimination
will also be influenced by the mineralogy of the ore concentrates.
Control of both the roasters and the reverberatory furnaces is
tocused on the combustion conditions in the equipment itself. Temperatures
at specific locations and visual observation of certain trays in the
hearth roaster are used to adjust air flows and hence combustion
Ln the roaster. In reverberatory furnaces, under established firing
c°nditions, oxygen analyzers at the furnace uptake and automatic draft
c°ntrol are used to achieve and maintain optimum control. Routinely
ayailable data have thus tended to be concentrated at these key metallurgical
°Perational steps. Conditions down the offgas system, with the exception
°* the operation of the waste heat boilers and the electrostatic precipitator
ave not warranted close attention in the past. With the advent of S02
Coiitrol legislation over recent years and the defined responsibilities
• shelter operators to limit emissions, there has arisen a need to know
6 S02 levels at the effluent stack; and the use of continuous S02
alyzers at the stack is becoming common.
179
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11.3 UNCERTAINTIES IN THE CHARACTERIZATION OF SMELTER WEAK
S02 GAS STREAMS
Effective matching of particular S0« control processes with the
weak SO, gas streams of individual copper smelters demands a thorough
understanding of the candidate control processes themselves, and a
detailed characterization of the subject SO,, gas streams. Unfortunate-
ly, the characterization of these gas streams has not resolved a number
of uncertainties related to the impact of operating practices and data
sources on gas flow variations and composition ranges.
As discussed above in Section 11.2, reverberatory gas streams are
subject to significant changes in both flow rate and S02 concentration.
For a paper study such as this current effort, the minimum information
needed to size an absorption-based system is 1) maximum gas flow rate
and 2) average S0» weight rate. The latter parameter must usually be
determined using the volume percent SO,, content of the gas and an
appropriate gas flow rate. Unless the values of SO,, content and gas
flow used for this calculation are approximately equal to the "average"
values of the system, the resulting SO,, weight rate can be substantially
in error, with distortion of both estimated costs and appropriate
control strategies.
Data available from individual smelters should be interpreted with-
in the framework of the conditions under which it is obtained or develop6**'
This has not always been possible. Some situations have been identified
where the gas stream characteristics are "composites" with the different
parameters having been determined at different locations in the system.
Routine data on sulfur trioxide levels, particulate loading after
the electrostatic precipitators, and the oxygen content of the gas
streams after the furnaces is generally not available simply because
such information has not been a necessary part of smelter operations.
It is also noted that although gas flow rates are determined on the
total gas volume including moisture, S02 determinations are usually on a
dry gas basis. In view of the uncertainties noted above, no correction
has been made for the SO,, content in calculating the weight rate of SO 2'
180
-------
11.4 UNCERTAINTIES IN THE APPLICATION OF ABSORPTION-BASED S02
CONTROL PROCESSES TO WEAK S02 STREAMS
In general, absorption-based SO^ control processes are in an early
stage of development, and there has been very little application to gas
streams with the characteristics provided by non-ferrous smelters.
While the development and pilot plant work conducted to date has been
mainly on utility plant flue gases, a greater understanding of the
Process relationships has resulted, and it is possible to identify at
least some of the areas of uncertainty these processes must accommodate
in the non-ferrous smelting field:
1) Variable S02 concentrations and gas volumes — Although most
of the candidate processes have demonstrated a measurable
degree of turn-down capability, the total effect of
relatively short cycle, wide variations in both S02
and gas flow in the absorption step is uncertain.
Liquid/gas (L/G) ratios can be established for the
maximum conditions but the changes in pH associated with
the changing conditions may pose operational problems,
particularly with the lime/limestone processes.
2) Oxygen levels — The weak S02 smelter gas streams are
invariably high in oxygen with values of up to 10-15
percent. This level of oxygen, together with the
possible presence of submicron-sized metallic
particulates acting as catalysts, offers the potential
for high rates of oxidation in the absorbent. High
oxidation rates increase purge requirements with the
loss of the active chemicals and inflate operating
expenses. The use of antioxidants may reduce
oxidation but they are not entirely effective and,
again, their use escalates operating costs significantly.
3) Utilization of S00 ~ The regenerable control processes,
with the appropriate auxiliary plant, produce sulfuric
acid, elemental sulfur, or liquid or gaseous sulfur
dioxide. Obviously, the choice of end product is an
economic one and will depend on internal requirements
for sulfuric acid or the marketing outlet for the
available options. In certain situations, the option
may be between acid production with neutralization
facilities to absorb over-market production, and the
production of elemental sulfur. It is beyond the scope
of this study to consider these aspects, but it is realized
that the total cost of a selected SO, control process for a
particular smelter could be substantially higher than
that cost which results from simply applying an auxiliary
sulfuric acid plant to the primary S02 control process.
181
-------
4) Adequacy of existing particulate removal facilities —
Recovery of solid material carryover from the roasters
and reverberatory furnaces is a routine practice in
copper smelters. In some cases, however, the electrostatic
precipitators in use are old, recovery efficiency is low,
and air infiltration is appreciable. The implementation
of an SCL control system would, in all likelihood, be
associated with replacement of the existing ESP. Where
such situations have been identified, the cost of a
replacement unit has been estimated but not directly
included in the cost of the S02 control system itself.
5) Cost uncertainties — Uncertainties in characterizing the
weak gas streams themselves will be reflected in the
associated cost structures but, at best, the cost estimates
developed for the specific smelters should be regarded as
order-of-magnitude only. Historically, as control methods
have progressed from pilot plant to demonstration scale
to full commercial application (for example the lime/
limestone system), estimated costs have escalated
dramatically. Estimated costs for the citrate process
appear to be very attractive today when viewed against
the costs for the other processes, but as this process,
too, continues through its development cycle, investment
costs may grow appreciably. The installation of new
processes in an on-line production facility also involves
significant cost uncertainties. Retrofitting costs and
the difficulties of scheduling, special safety requirements,
provision of temporary production facilities, etc., are
all variables with uncertain limits. Allowances have been
made in the smelter application costs, but this uncertainty
should be noted. The cost structures for each of the
basic control systems have been escalated to a mid-1974
base, but cost escalation levels recently experienced make
it difficult to project future costs. This situation
should certainly be kept in mind in evaluating the developed
costs in the subsequent sections.
182
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12.0 APPLICATION OF ABSORPTION-BASED S02 CONTROL
SYSTEMS TO THE WEAK S02 GAS STREAMS OF
U.S. PRIMARY COPPER SMELTERS
12.1 THE ANACONDA COMPANY-ANACONDA, MONTANA
The Anaconda Smelter with a nominal capacity of 2000 TPD of con-
centrate presently operates with 4 reverberatory furnaces and 4 on-line
Pierce-Smith converters. Reverberatory gases from one furnace pass
through a waste heat boiler while the gases from the other three units
are water quenched in a cooling chamber. The reverberatory gases are
then mixed with a portion of the converter offgases before passing
through an electrostatic precipitator and being vented up a 925 foot
stack. About 83,000 SCFM of the converter gases are directed to a
double absorption sulfuric acid plant presently producing about 400 TPD
°f acid.
The company is presently in the final stages of a major project to
convert its smelting operation from reverberatory furnaces to a fluo-
solids roasting and electric furnace process. S02 content of the off-
gases is expected to be around 6.5 percent and the present acid plant
will be upgraded to 1100 TPD acid to accommodate this loading. Startup
is scheduled for October 1975 with full on-line operation expected by
the end of the year.
In view of this situation, no absorption-based control systems will
be considered.
183
-------
184
-------
12.2 ASARCO — EL PASO SMELTER, TEXAS
12.2.1 Smelter Characteristics
This smelter began operations in 1905. It is a custom smelter with
a nominal capacity of around 1000 tons of concentrate per day. The
concentrates are treated in 4 multi-hearth roasters, smeltered in a
single reverberatory furnace and 2 Fierce-Smith converters. The roaster
gases pass through a settling flue and join the reverberatory gases
Prior to a spray chamber and the electrostatic precipitator. Presently
in the planning stage, with a 1978 target date, is a proposal to treat
the roaster gases together with the gases from a lead smelter sintering
machine in a new 500 TPD sulfuric acid plant.
The reverberatory gases pass through waste heat boilers and mix
with the roaster gases; the total gas stream, after cooling in a spray
chamber, is treated in an ESP before being vented up an 825 ft stack.
The offgases from the 2 operating converters (a third unit is on
standby) pass through waste heat boilers, spray chamber, and an electro-
8tatic precipitator before feeding a double contact, 525 TPD sulfuric
a°id plant rated at 60,000 SCFM at 4.5 percent S02, commissioned in
1972.
The El Paso smelter is located close to the city of El Paso on a
fairly congested plant site. Steam and power facilities are fully
l°aded. Water is available.
^•2.2 Weak SO., Streams
Although there are two sources of weak SC^ streams in this smelter,
the mixing of these streams prior to the electrostatic precipitators
tesults in one combined dilute stream being vented to atmosphere. A
fl°w schematic is provided in Figure 12.2-1 and indicates the individual
Stream parameters which have been developed to represent this smelter's
°peration. It is noted that these data represent El Paso's operation
°nly under the conditions of processing a specific ore concentrate.
Slnce this smelter is a custom smelter, ore concentrate characteristics
and hence operating conditions and S02 elimination may change markedly.
he reverberatory fuel/charge ratio for the reported period was low and
Qf\A
8 flot necessarily represent the long range conditions.
185
-------
On the basis of the data provided for the El Paso operation, the
S09 content of the reverberatory gases range from 0.06 to a high of 0.78
percent over a cycle of 3 chargings. At the normal operating rate of a
charge taking place approximately every 20 minutes, the related S02
level swings make the determination of a meaningful SC^ level for S02
control sizing extremely difficult. The practice of "soot" blowing the
waste heat boilers approximately every hour also contributes to the gas
volume and SO- content variability. The "average" figures provided must
be taken as indications only. The S02 levels for the El Paso multiple
hearth roasters also show wide variations over a 2 hour cycle, and an
approximate "average" has been estimated.
The parameters of the combined stream flow presently being vented
represent "average" reported gas flows, with the SO^ and oxygen contents
being estimated for the individual stream information.
The weak gas streams of the ASARCO El Paso smelter under present
operation are thus characterized as shown in Table 12.2-1.
12.2.3 S£2 Control Process Selection
As noted in Section 11, custom smelters pose a large degree of
uncertainty when it comes to selecting and sizing S0~ control processes.
The "average" characteristics identified for the El Paso Smelter provide
a particularly doubtful basis. Although this smelter is nominally
equivalent to the smelting capacity of ASARCO's Tacoma smelter, the
particular conditions under which this smelter was operating and for
which the reported data applies were very different, and the resulting
S02 levels are markedly different from those of Tacoma. Since it is
quite feasible for the El Paso smelter to handle concentrates similar to
those being handled at Tacoma, or vice versa, there is some validity in
considering these two smelters as having the same parameters as far as
selecting and sizing S02 control processes. That is, the S02~handling
section of the process should be sized for the maximum average S02
levels expected. This approach will be taken for the combined roasters
and reverberatory streams. A secondary situation will be considered
where the roaster gases are already treated in the planned second acid
plant and only the reverberatory gas stream at the maximum expected S02
concentration and gas flows will require scrubbing treatment.
186
-------
— EL PASO, TEXAS
SMELTER FLOW SCHEMATIC
95-J 10,000 SCFM
0.7-1.4% SO2
(AV = 0.85%)
330° F
ROASTERS
(4)
60-80,000 SCFM
0.06-0.78% S00
oo
(
REVERB
(1)
02 J
V(AV = 0.35%) J
650°F y
\A/MR
SPRAY
CHAMBER
ESP
^ STACK
825 FT
------- 1
, CONVERTERS
'
I 1
-•} WHB {—
I I
| SPRAY '
I CHAMBER |
I ]
•j ESP f
I j'
r
1 ACID
J PLANT
I (D.C.)
I
I
j > STACK
I
(2 OPERATE AT
TIME)
525 TPD
[60,000 SCFM @ 4.5% SO2]
FIGURE 12.2-1
-------
Table 12.2-1. ASARCO - EL PASO SMELTER, TEXAS
CHARACTERIZATION OF WEAK S02 GAS STREAMS
BASED ON PRESENT OPERATION.
Stream
Roasters (4)
Reverbs. (1)
(after W.H.B.)
Combined streams
after ESP<3>
Volume
SCFM
95-110,000
60-80,000
190-220,000
Av. S02 Content
Sizing basis (Ibs/hr)
8,600
2,500
11,100
Equivalent Av.
SO Content (%)
0.85
0.35
0.52
% Oxygen^2)
5-7
4-6
7-9
Temp. °F
330
650
250
00
oo
Notes. (1) No data on SO or particulate loading available.
(2) Oxygen content of gas streams estimated.
(3) Additional ventilation gas streams are vented into the
flue system, but since they carry little or no S02>
they have not been included in the combined roaster
and reverberatory stream values. The provided values
do include the additional air infiltration after
the W,H.B.
-------
The "adjusted" parameters for the El Paso smelter will thus be
defined as shown in Table 12.2-2.
Although the S02 absorption section of any control system is largely
sized by the gas flow rate, the liquid/gas ratios are related to the SO
concentration and would be fixed at the higher levels. To provide a
reasonable sizing basis for the S02-handling section, additional surge
tanks would be necessary after the absorption system to give an aver-
aged S02~content absorbent input.
Of the throwaway processes, only the double alkali process appears
to be practical. The high peaking concentration of S02 would necessi-
tate 2-stage absorption for the limestone process and on the basis of
Past experience, the fluctuating S02 content may lead to pH control
Problems with concomitant scaling and operating problems.
The regenerable processes are all theoretically applicable although
he relatively high oxygen levels present in the blended roaster/rever-
eratory gases pose major problems in terms of economics and operation.
TVi
ne two control processes which are the least affected by oxidation are
citrate process and the magnesium oxide process. The resulting,
rather weak (15%) S09 obtained from the regeneration process of the
2
MgO
Process could not be handled in the existing acid plant and an additional
cid plant with full gas cleaning facilities would be necessary. Although
me oxidation would still be experienced with the citrate process, the
ch lower level suggests that this process with its end product of
Cental sulfur offers practical advantages. It is recognized, how-
eri that the status of this process—only limited pilot plant experience-
fl*» J
the requirement for natural gas do constitute uncertainties.
The water quenching of the roaster/reverberatory gas stream and the
e,.
crate temperature which exists after the electrostatic precipitator
Approximately 250°F) will result in very little additional gas cooling
bgjn
"8 required before the S02 absorption section.
1>P_2 Contro1 Process Costs
Conditions specific to this smelter which will directly affect the
tal and operating costs of the available S02 control processes
nclude the following:
189
-------
1) The mixing of the roaster and reverberatory gases and
quenching them in a spray chamber provide both cooling
and some degree of cleaning. The temperature of this gas
stream after the electrostatic precipitators is approximately
250°F, and only an additional water quench to bring the
temperature down to about 130°F should be necessary before
the SO- absorption column.
2) Retrofitting costs will reflect the difficulties of making
the necessary tie-ins with appropriate bypass and dampers to
the rather short (approximately 100 ft) flue system existing
between the ESP and the stack. A factor of 50 percent of
the capital cost of the prequench and SO- absorption system
has been assumed.
3) The variability of both gas flow and S02 content in the
reverberatory gas stream will require the provision of larger
surge capacity after the absorption system to provide a uniform
SO. rate to the regeneration section, and an additional 5
percent of the capital investment of the SO- absorption
section has been provided to cover this provision.
4) Additional service facilities such as substations, power
distribution, and a water treatment plant to support a scrubbing
process will be necessary. Estimated costs have been based on
the unit values established for each control process.
5) The implementation of a major process change in an operating
plant incurs substantial additional costs not readily
identifiable in advance, but nevertheless directly affecting
the total capital cost of the project. An estimate of 20
percent of the total boundary limit costs for each process
has been allowed.
With the weak SO- stream parameters of this smelter adjusted to those of
ASARCO's Tacoma smelter, it is apparent that SO- control costs will
parallel those of Tacoma. However, there are certain differences which
will have a small effect on final investment costs. Retrofitting the
absorption system itself in the El Paso smelter should not pose the same
difficulties as might be expected in the Tacoma smelter. Special treat-
ment of recovered ESP particulates for secondary metal recovery with the
associated auxiliary gas stream being vented into the roaster/rever-
beratory flue system as presently practiced at Tacoma is not now part of
190
-------
The El Paso operation, and should such an operation become necessary at
El Paso through ore concentrate changes, the capital costs associated
with providing auxiliary gas cleaning are independent of the S02 control
system. Hence no provision of auxiliary gas cleaning facilities to
reduce the total gas volume through an SC^ control system is necessary.
Summaries of capital and annual operating costs for selected S02
control systems on the combined roaster and reverberatory streams of the
El Paso smelter are provided in Tables 12.2-3 and 12.2-4, respectively.
Summaries of capital and operating costs for selected S02 control systems
applied to the reverberatory gas stream only are provided in Tables
12.2-5 and 12.2-6, respectively. Individual calculation sheets for the
applied processes—namely, double alkali, magnesium oxide and citrate process-
are provided under the appropriate smelter heading in Appendix G.
191
-------
Table 12.2-2. ASARCO - EL PASO SMELTER, TEXAS.
CHARACTERIZATION OF WEAK S02 GAS STREAMS
ADJUSTED FOR CONCENTRATE CHANGE.
Stream
Roasters (4)
Reverbs . (1)
(after W.H.B.)
Combined streams
after ESP<3'
Volume
SCFM
95-110,000
70-100,000
200-250,000
Av. S02 Content
Sizing basis (Ibs/hr)
9,200
13,800
23,000
Equivalent Av.
S02 Content (%)
0.9
1.6
1.0
(2)
% Oxygen* '
5-8
5-8
8-12
Temp. °F
330
600
250
H1
VO
Notes.
(1) No data on SO or particulate loading available.
(2) Oxygen content of gas streams estimated.
(3) Additional ventilation gas streams are vented into the
flue system, but since they carry little or no S02>
they have not been included in the combined roaster
and reverberatory stream values. The provided values
do include the additional air infiltration after
the W.H.B.
-------
Table 12.2-3. ASARCO - EL PASO, TEXAS
CAPITAL COSTS FOR SO2 CONTROL PROCESSES ON COMBINED ROASTER AND REVERBERATORY FURNACE OFF-GASES
ITEM
1. Boundary limits retrofitted
fitted primary system
2. Auxiliary Plant:
(a) H SO plant with dry
gas cleaning
TOTAL Boundary Limit Cost
3. Total Service Costs
4. Associated Investment Costs
(a) new ESP
(b) General site- related
costs
Total Capital Investment
Capital cost/annual ton S02
Removed
Primary S0_ Control Method
Double Alkali
(95% Removal)
10,390,000
N/A
10,390,000
550,000
2,080,000
13,020,000
$146
Magnesium Oxide
(90% Removal)
9,270,000
4,350,000
13,620,000
570,000
2,720,000
16,910,000
$206
Citrate
(95% Removal)
10,390,000
N/A
10,390,000
800,000
2,080,000
13,270,000
$149
COMMENTS
VO
U>
-------
Table 12.2-4. ASARCO - EL PASO, TEXAS
ANNUAL OPERATING COSTS AND TOTAL NET ANNUALIZED COSTS FOR S02 CONTROL PROCESSES ON
COMBINED ROASTER AND REVERBERATOR? FURNACE OFF-GASES
ITEM
V
1. Gas conditioning and S02
absorption
2. S02 handling
3. Labor
4. Maintenance, Ins. & Taxes
5. Total Annual Operating Cost
6. Auxiliary Plant Total
Operating Costs
(a) Sulfuric acid plant
7. Total Annual Operating Cost
8. Annualized Capital Cost
9. Total Net Annualized Cost
10. Cost/year/ton S02 removed
11. Cost/lb copper^1)
Primary S0_ Control Method
Double Alkali
(95% Removal)
179,000
3,906,000
158,000
742,000
4,985,000
N/A
4,985,000
1,712,000
3,888,000
$44
3.35c/lb
Magnesium Oxide
(90% Removal)
179,000
2,542,000
262,000
778,000
3,761,000
520,000
4,281,000
2,224,000
3,909,000
$48
3.370/lb
Citrate
(95% Removal)
179,000
2,226,000
280,000
748,000
3,433,000
N/A
3,433,000
1,745,000
3,106,000
$35
2.68c/lb
COMMENTS
Based on smelter rated capacity of 1000 TPC concentrate , 18% copper, operating 340 days/year.
Annual production 58,000 TPY. Zero net-back to smelter for acid or sulfur produced.
-------
Table 12.2-5. ASARCO - EL PASO, TEXAS
CAPITAL COSTS FOR SO2 CONTROL PROCESSES ON REVERBERATORY FURNACE OFF-GASES ONLY
ITEM
1. Boundary limits retrofitted
primary system
2. Auxiliary Plant:
(a) H2S04 plant with dry
gas cleaning
TOTAL Boundary Limit Cost
3. Total Service Costs
4. Associated Investment Costs
(a) new ESP
(b) General site-related
costs
Total Capital Investment
Capital cost/annual ton S02
Removed
Primary S0_ Control Method
Double Alkali
(95% Removal)
7,280,000
N/A
7,280,000
420,000
1,460,000
9,160,000
$172
Magnesium Oxide
(90% Removal)
6,350,000
3,100,000
9,450,000
420,000
1,890,000
11,760,000
$240
Citrate
(95% Removal)
7,580,000
N/A
7,580,000
600,000
1,520,000
9,700,000
$182
COMMENTS
vo
Ul
-------
Table 12.2-6. ASARCO - EL PASO, TEXAS
ANNUAL OPERATING COSTS AND TOTAL NET ANNUALIZED COSTS FOR SO CONTROL PROCESSES ON
REVERBERATORY FURNACE OFF-GASES ONLY
ITEM
1. Gas conditioning and SO.,
absorption
2. S02 handling
3. Labor
4. Maintenance, Ins. & Taxes
5. Total Annual Operating Cost
6. Auxiliary Plant Total
Operating Costs
(a) Sulfiric acid plant
7. Total Annual Operating Cost
8. Annualized Capital Cost
9. Total Net Annualized Cost
10. Cost/year/ton S0_ removed
11. Cost/lb copper t1'
Primary S02 Control Method
Double Alkali
(95% Removal)
82,000
2,340,000
158,000
520,000
3,100,000
N/A
3,100,000
1,205,000
2,524,000
$47
2.18c/lb
Magnesium Oxide
(90% Removal)
82,000
1,523,000
260,000
535,000
2,400,000
370,000
2,770,000
1,546,000
2,611,000
$52
2.25c/lb
Citrate
(95% Removal)
82,000
1,337,000
280,000
546,000
2,245,000
N/A
2,245,000
1,276,000
2,133,000
$40
1.84c/lb
COMMENTS
V0
Based on smelter rated capacity of 1000 TPD concentrate, 18^ copper, operating 340 days/year.
Annual production 58,000 TPY. Zero net-back to smelter for acid or sulfur produced.
-------
12.3 ASARCO — HAYDEN SMELTER, ARIZONA
12.3.1 Smelter Characteristics
This custom smelter with a nominal capacity of 1800-2000 tons of
concentrate per day began operations in 1912. There are 12 multiple
hearth roasters with normally 6-8 in operation at any one time. Gases
from the roasters enter a settling chamber where they are blended with
the reverberatory furnace gases.
Roasted calcine is handled in two reverberatories—one 27 feet wide
and 112 feet long, and the other 23 feet wide and 105 feet long. Charging
is accomplished using Wagstaff guns. Gases pass through waste heat
Boilers and are cooled in a spray chamber before passing into the settling
chamber to be blended with the roaster gases. The combined stream is
treated in an electrostatic precipitator for particulate recovery
before being vented to a new 1000 ft stack completed in 1974.
There are five Fierce-Smith converters provided with water cooled
hoods, with normally three being in operation at a time. The converter
Sases pass through cyclones, a water spray chamber, and an electrostatic
Precipitator before passing to the gas conditioning section of a single
c°ntact sulfuric acid plant rated at 750 TPD treating approximately
OOO SCFM of 4.0 percent SO™ content feed. The tail gas is vented up
1000 ft stack. Some availability and efficiency problems have been
experienced with this acid plant, but recent maintenance programs may
corrected this situation.
The smelter is located at about 2000 ft elevation, adjacent to the
Copper Corporation's Hayden Smelter site. The surrounding
which is available to the smelter for industrial purposes is limited.
Uch of it is under lease as grazing land or is Federally-owned land.
The smelter fully utilizes its present steam generating capability
°r power generation, and available power is fully loaded. Water avail-
a°ility is good although it is hard and typical of Arizona water supplies.
197
-------
12.3.2 Weak SO,, Streams
As with the other ASARCO copper smelters, both the roaster gases
and the reverberatory gases are weak SCL streams and these are mixed
prior to the electrostatic precipitator and then vented as a combined
stream. The flow schematic provided in Figure 12.3-1 indicates the
individual stream parameters based on Hayden's operation under the
conditions of processing a particular charge of ore concentrates.
Again, as noted with ASARCO's other copper smelters, ore concentrate
characteristics may change with corresponding changes in operating
conditions and SO- elimination.
Data from the Hayden smelter indicates that SO- levels in the
combined gases from the two reverberatory furnaces, as determined after
the waste heater boilers, respond markedly to the furnace charging se-
quences. At the normal operating rates, there are about 16 loads per
hour smelted in the larger reverberatory, and 10 loads per hour smelted
in the smaller furnace. Within this period the determined S0~ level
appears to vary from a low of approximately 0.1 percent to a peak value
of 1.4 to 1.7 percent as charging takes place. An "average" S02 value
of 0.5 percent has been estimated for this combined reverberatory gas
stream after the spray tower.
The multiple hearth roaster gases also show variability in S0»
content and the effect of appreciable air dilution. The gas volume has
been "normalized" and an approximate "average" value for S02 content
estimated.
The characteristics of the combined stream flow after the electro-
static precipitator have been developed from the individual stream
information with appropriate allowance for additional air dilution.
The weak gas streams of the ASARCO Hayden Smelter have been charac-
terized as indicated in Table 12.3-1.
12.3.3 S00 Control Process Selection
^ —— —_____^_^__^_______^____
The S02 control processes for this smelter will be applied to the
blended roaster and reverberatory stream after the electrostatic pre-
cipitator. The selection and sizing of specific processes will be based
on the parameters defined in Table 12.3-1.
198
-------
ASARCO — HA YDEN, ARIZONA
SMELTER FLOW SCHEMATIC
140-150,000 SCFM
1.05-1.65% SO2
-------
The constraint on available land around the Hayden Smelter for
sludge disposal ponds and the variable SCL level of the effluent gases
appear to preclude application of the lime/limestone process. The
double alkali throwaway process will be susceptible to the relatively
high oxygen levels in these streams, but the regeneration chemistry of
this system suggests that the process may offer potential in this appli-
cation. The resulting filtered sludge would have to be trucked from the
smelter site, but for the purposes of this study, it is assumed that a
suitable landfill location is available within 10 to 15 miles of the
smelter.
Oxidation also imposes significant penalties on all the regenerable
processes with perhaps the exception of the citrate process and the
magnesium oxide process. The sulfite and bisulfite products usually
resulting from an S02 absorption process are oxidized to non-regenerable
sulfates which must then be removed from the system. These purge require-
ments result in the loss of the active absorbent chemicals or, alternative*
ly, the treatment of these heavier purge streams will incur significant
additional capital investment and increased operating expenses.
The citrate process with its tolerance for relatively high oxygen
levels without an appreciable increase in S02 oxidation levels, appears
to offer the outstanding potential with an end product of elemental
sulfur. In the magnesium oxide process sulfite oxidation appears to be
self-limiting at around 15 percent MgSO,, but if this situation does not
exist, purging and secondary removal of the MgSO, will be necessary; or
alternatively, if the MgSO^ concentration builds up to approximately 33
weight percent, MgS04«7H20 will be precipitated with the MgS03«6H20 and
will be reduced during the calcination step. This process requires the
provision of additional gas cleaning facilities to handle the approximate*"
15 percent SO, regenerated stream and an auxiliary sulfuric acid plant-
^* %
The direct application of a sulfuric acid plant to gas streams of this
volume and with such low levels of S00 would demand unrealistically i
tic'
sized preheating and refrigeration facilities and is obviously an imprflC
approach.
The S02 control processes which appear to offer reasonable potenti*1
for application in this smelter are the double alkali, citrate, and
magnesium oxide processes.
200
-------
12.3.4 SOp Control Process Costs
The basic cost structures for the selected processes for this
smelter have been modified to take into consideration the special factors
and conditions associated with this operation.
1) The practice common to the ASARCO copper smelters of mixing
the roaster and reverberatory gases and quenching the
combined stream in a spray chamber provides both cooling
and a degree of cleaning. The temperature of this gas stream
after the electrostatic precipitator is between 250 and 270°F
and only an additional water quench to bring the temperature
down to about 130°F should be necessary before the absorption
column. An appropriate cost allowance has been made for
this pre-absorption quench.
2) Retrofitting costs to accommodate this S02 absorption
column and prequench with the necessary damper system
have been allowed for on the basis of 50 percent of the
capital cost of the S02 absorption column and prequench.
3) An additional 5 percent of the capital investment of the
S09 absorption section has been provided to allow for
increased surge capacity after the absorption section
to accommodate the variable S02 content and gas flow rates,
and to provide a uniform S02 rate to the regneration section.
4) Additional power and water treatment facilities will be
required to support a scrubbing process, and costs have
been based on the unit relationships established for each
control process and the cost-capacity relationships
provided in Appendix F.
5) A general site cost factor of 20 percent of the total
boundary cost for each process has been allowed to cover
those additional but not readily identifiable charges
associated with the implementation of a major process
change in an operating plant.
The capital and operating cost structures for the three processes
Elected as applicable for the Hayden Smelter—double alkali, citrate,
and magnesium oxide processes—are provided in Tables 12.3-2 and 12.3-3,
tesPectively. Individual computation sheets for each process are pro-
bed under the appropriate smelter name in Appendix G.
201
-------
Table 12.3-1. ASARCO - HAYDEN SMELTER, ARIZONA.
CHARACTERIZATION OF WEAK S02 GAS STREAMS.
Stream
Roasters (6-8)
Reverbs. (2)
(after spray
chamber)
Combined streams
(after ESP's)
Volume
SCFM
140-150,000
150-180,000
350-400,000
Av. S02 Content ^
Sizing basis (Ibs/hr)
18,400
8,100
26,500
Equivalent Av.
S02 Content (%)
1.25
0.5
0.7
% Oxygen(2)
4-8
4-8 :
8-12
Temp. °F
250
330
270
N»
O
N>
Notes. (1) No data on SO- or particulate loading available.
(2) Oxygen content of gas streams estimated.
-------
Table 12.3-2. ASARCO - HAYDEN, ARIZONA
CAPITAL COSTS FOR SO2 CONTROL PROCESSES ON COMBINED ROASTER AND REVERBERATORY FURNACE OFF-GASES
ITEM
1. Boundary limits retrofitted
primary system
2. Auxiliary Plant:
(a) H2S04 plant with dry
gas cleaning
TOTAL Boundary Limit Cost
3. Total Service Costs
4. Associated Investment Costs
(a) new ESP
(b) General site-related
costs
Total Capital Investment
Capital cost/annual ton S09
Removed
Primary S02 Control Method
Double Alkali
(95% Removal)
11,940,000
N/A
$11,940,000
580,000
$ 2,390,000
$14,910,000
$146
Magnesium Oxide
(90% Removal)
9,990,000
4,750,000
$14,740,000
610 ,-000
$ 2,950,000
$18,300,000
$194
Citrate
(95% Removal)
11,990,000
N/A
$11,990,000
830,000
$ 2,400,000
$15,220,000
$149
COMMENTS
-------
Table 12.3-3. ASARCO - HAYDEN, ARIZONA
ANNUAL OPERATING COSTS AND TOTAL NET ANNUALIZED COSTS
ITEM
1.
2.
3.
4.
5.
6.
7.
8.
9.
10.
11.
Gas conditioning and S02
absorption
S02 handling
Labor
Maintenance, Ins. & Taxes
Total Annual Operating
Cost
Auxiliary Plant total
Operating Costs
(a) Sulfuric acid plant
Total Annual Operating
Cost
Annualized Capital Cost
Total Net Annualized Cost
Cost/year/ton S02 removed
Cost/lb copper (•*-'
Primary S0_ Control Method
Double Alkali
(95% Removal)
270,000
4,479,000
158,000
851,000
$5,758,000
N/A
$5,758,000
$1,961,000
$4,478,000
$44
1.93
-------
12.4 ASARCO — TACOMA, WASHINGTON SMELTER
12.4.1 Smelter Characteristics
This smelter handles custom smelting exclusively. Of the 10 multiple
Dearth roasters available, 4-6 are usually on-line at a time. They are
°ld with high leakage rates and correspondingly high dilution of the
r°asting off gases. These gases are combined with the reverberatory
furnace gases in a spray chamber prior to the electrostatic precipitators.
r gases from the arsenic recovery plant, ventilation gases, etc.,
c°ntaining little or no S02 are also mixed with the roaster and rever-
eratory gases before the electrostatic precipitator.
One on-line reverberatory furnace handles the roaster calcines with
c arging normally taking place 4-5 times/hour. The off gases pass through
aste heat boilers before being combined with the roaster gases and
rther cooled in a spray chamber. They are then mixed with other
elter gases and the combined stream passes through two electrostatic
6cipitators in series before being vented up a 535 ft stack. Two oil
°ers are used at the base of the stack to provide adequate buoyancy.
The offgases from the two on-line converters are subjected to a
Prehensive gas conditioning and cleanup sequence including electro-
tic precipitators, scrubbers and mist ESP's before feeding a Monsanto
TPD sulfuric acid plant and a dimethylaniline scrubbing system to
°
-------
will reflect this situation in both gas volumes and SO,, content. The
volume of exhaust gases is influenced by the firing rate or fuel/charge
ratio and this in turn is related to the "smelterability" of the charge
ore material. As discussed in Section 11, fluctuations of the SQ^
content in the exhaust gases from both the roasters and reverberatory
furnace are normal conditions. On the basis of the data provided by
ASARCO, the S02 content of the Tacoma reverberatory offgas sampled after
the W.H.B. undergoes wide swings of as much as 4.8 percent down to 0.5
percent during the charging sequence. At a rate of 4-5 charges per hour
(2 larry-cars/charge), it is apparent that the S02 level and the volume
of gas will be continually varying over relatively short time periods.
As a basis for sizing the sulfur dioxide-dependent section of any control
process, i.e., the regeneration section or the raw material requirements
of the throwaway processes, an average S02 level of approximately 1.7
percent can be estimated. Because of the variability associated with
custom smelting, even this value could change significantly (see data
for the ASARCO, El Paso and Hayden Smelters).
No direct data for the Tacoma roasters are available but on the
basis that smelting capacity at Tacoma is approximately equivalent to
that of the ASARCO El Paso Smelter, the total roaster gas volume and S02
concentration noted for this latter smelter have been used for defining
the Tacoma system. The SO,, concentration of 0.9 percent represents an
approximate "average" of the concentration profile available.
The present weak gas streams in the Tacoma Smelter are thus character*-2
/
as shown in Table 12.4-1.
12.4.3 Selection of SO^ Control Processes
Under the present operational mode of the Tacoma smelter, the
combined roaster and reverberatory gas streams after the electrostatic
precipitators include an appreciable quantity of essentially "zero-S02"
gases from auxiliary operations conducted at the Tacoma site. It has
been estimated that these gases could amount to as much as 30,000-40,000
SCFM. In considering the application of an S0? control process to the
combined roaster and reverberatory, it is apparent that including these
auxiliary gas streams as part of the total scrubbing input gas load
results in a significantly bigger scrubbing system with higher capital
and operating costs. On the other hand, if these gas streams are
206
-------
ASARCO — TACOMA, WASHINGTON
SMELTER FLOW SCHEMATIC
9 5-110,000 SCFM
0.7-4.8% SO2
(AV = 0.9%)
330°F
ROASTERS
(8)
(6 IN OPERATION
AT TIME)
70-100.000 SCFM
0.7-^.8% SO2
(AV = 1.7%)
NJ
O
REVERB
(1)
GASES FROM
AUXILIARY
OPERATIONS
STACK
565 FT
I
I 1
[ CONVERTERS [_
I (4) I
I 1
|{2 IN OPERATION]
[_ AT TIME) J
H2S04
GAS CLEANING J
AND r
CONDITIONING |
1
1 PLANT [
I 150 TPD I
I 1
r 1
I DMA I
, PLANT j~~
I 1
STACKS
FIGURE 12.4-1
-------
diverted from the roaster/reverberatory electrostatic precipitators,
alternative processing facilities must be provided, and the capital cost
and operating cost of such facilities should be charged against the S02
control process applied.
For the purposes of this study, it will be assumed that alternative
processing will be provided for the auxiliary gases and that the S02
control processes will be sized for the combined roaster/reverberatory
stream as characterized in Table 12.4-1.
The urbanized location of this smelter- with its limited land avail-
ability precludes the application of the lime/limestone process. Although
problems might be anticipated with the routine trucking of sludge wastes
from the double alkali process through an urban area, there is a possi-
bility of using this material with proper fixation techniques as fill
material in reclaiming land from the sea area.
The highly variable S02 content of the reverberatory stream will
impose additional cost penalities on scrubbing processes by virtue of
the higher liquid/gas (L/G) ratios required to handle the weak S02
levels and, in the case of the xylidine and ammonia processes, larger
heat exchange units for the inter-tray coolers. The oxygen level of the
combined streams has been estimated at around 8-12 percent, and at this
level, it is expected that the Wellman-Lord, xylidine and ammonia pro-
cesses all would be adversely affected in both capital and operating
costs by the higher oxidation rates. The direct use of a sulfuric acid
plant on this weak stream is clearly an unattractive alternative when
assessed in economic terms and energy requirements.
The three processes which appear to have the capability of accommo'
dating the oxidation problem and the variable SCL level with the least
penalties are the double alkali process, the citrate process and the
magnesium oxide process, although it is noted that the energy require-
ments of the latter process together with the requirements for gas
conditioning facilities and an auxiliary sulfuric acid plant pose a
substantial economic, premium. It is noted that as a result of the in-
process quenching of these gas streams and the low temperature-of the
final mixed gases (250°F), little if any separate gas cooling and
conditioning will be required for r.he selected control processes.
208
-------
12.4.4 SOo Control Process Coses
The capital and the operating costs developed for the control
Processes judged as applicable to the Tacoma smelter recognize a number
°f conditions which are peculiar to this smelter.
1) The blending of the roaster and reverberatory gases and
subjecting them to a water quenching step provide both
cooling and conditioning prior to the electrostatic
precipitators. The gases exiting the ESP's are at a
temperature of approximately 250°F and an additional
simple quench to approximately 130°F prior to the S02
absorption column should be adequate. Any particulate
carryover would be removed from the S02 control process
itself via the purge system. Operating costs will be
incrementally increased by approximately 80 gal/min of
treated water.
2) Retrofitting costs—i.e., making the necessary tie-ins to
the existing ductwork after the ESP's, with appropriate
bypass and dampers for the S02 absorption column—would be
expected to be high because or the age of the smelter and
the congested space. A retrofit factor of 75 percent of
the cost of the prequench and S02 absorption system has been
assumed.
3) To accommodate the variability of both gas flow and SO,
content in these gas streams, larger surge capacity after
the absorption section will be necessary to provide a
uniform S02 rate to the regeneration section. An additional
5 percent of the capital investment of the S02 absorption
section has been provided.
4) The auxiliary gases which are presently vented into the gas
handling system prior to the electrostatic precipitators
contain little or no S02 and the approach has been taken to
eliminate these streams (estimated at approximately 40,000
SCFM) from the gas flow to the S02 control system. However,
a replacement particulate removal system would have to be
provided and the cost should be allocated against the total
cost for S02 control. A bag filter system has been selected
and costed accordingly.
5) The capital costs of service facilities such as substations,
power distribution, and water treatment plants have been
estimated based on the unit values established for each
specific control process and the cost data provided in Appendix
F. Process water has been assumed to need a degree of chemical
treatment.
6) The implementation of a major process change in a congested
operating plant incurs substantial additional costs not
readily identifiable in advance but nevertheless directly
affecting the total capital cost of the project. An
estimate of 20 percent of the total boundary costs for
209
-------
Individual calculation sheets for the applied processes are provided
under the appropriate smelter heading in Appendix G and a summary of
these capital costs for the Tacoma smelter is provided in Table 12.4-2.
Associated annual operating and total annualized costs are provided in
Table 12.4-3.
210
-------
Table 12.4-1. ASARCO - TACOMA SMELTER, WASHINGTON
CHARACTERIZATION OF WEAK S<>2 GAS STREAMS
Stream
Roasters (6)
Reverb (1)
(after W.H.B.)
Combined Streams
after ESP's3
Volumes
SCFM
95-110,000
70-100,000
200-250,000
Av- S0? Content
Sizing Basis (Ibs'/hr)
9,200
13,800
23,000
Equivalent Av.
S02 Content (%)
0.9
1.6
1.0
% Oxygen
5-8
5-8
8-12
Temp°F
330°
600°
250°
NOTES:
1) No data an SO- or particulate loading available.
2) Oxygen content of gas streams estimated.
3) At the Tacoma smelter additional gas streams estimated at 30-50,000 SCFM are directed into the
system prior to the ESP's. These streams carry little or no S02 and they have not been
included in these combined roaster and reverberatory stream values. There is additional air
infiltration after the W.H.B.
-------
Table 12.4-2. ASARCO - TACOMA, WASHINGTON
CAPITAL COSTS FOR S02 CONTROL PROCESSES ON COMBINED ROASTER AND REVERBERATORY FURNACE OFF-GASES
ITEM
1. Boundary limits retrofitted
primary system
2. Auxiliary Plant:
(a) H-SO, plant with dry
gas cleaning
TOTAL Boundary Limit Cost
3. Total Service Costs
4. Associated Investment Costs
(a) Bag filter system
for auxiliary gases
(b) General site-related
costs
Total Capital Investment
Capital cost/annual ton S02
Removed
Primary SO- Control Method
Double Alkali
(95% Removal)
11,000,000
N/A
11,000,000
550,000
600,000
2,200,000
$14,350,000
$161
Magnesium Oxide
(90% Removal)
9,960,000
4,350,000
14,310,000
570,000
600,000
2,860,000
$18,340,000
$224
Citrate
(95% Removal)
11,000,000
N/A
11,000,000
800,000
600,000
2,200,000
$14,600,000
$164
COMMENTS
No gas conditioning provided
but cost allowance for pre-
absorber quenching to 130°F.
Retrofit allowance 75% absorption
section.
Gas cleaning provided by high
temperature bag filters rather
than wet scrubbing.
Bag filter systems for auxiliary
gases previously vented through
ESP's.
Related to difficulties of
scheduling, interruptions, etc.,
in operating plants.
ro
-------
1.2. 4-3. ASARCO - TACOMA, ffASHIffGTOfT
ANNUAL OPERATING COSTS AND TOTAL NET ANNUALIZED COSTS FOR SO- CONTROL
PROCESSES ON COMBINED ROASTER AND REVERBERATORY FURNACE OFF-GASES
ro
ITEM
1. Gas conditioning and S02
absorption
2. S02 handling
3. Labor
4. Maintenance , Ins . & Taxes
5. Total Annual Operating Cost
6. Auxiliary Plant Total
Operating Costs
(a) Sulfuric acid plant
7. Total Annual Operating Cost
8. Annualized Capital Cost
9. Total Net Annualized Cost
10. Cost/year/ton S0» removed
11. Cost/lb copper (1'
Primary S0_ Control Method
Double Alkali
(95% Removal)
179,000
3,906,000
158,000
799,000
5,042,000
N/A
$5,042,000
$1,887,000
$4,050,000
$46
2.28e/lb
Magnesium Oxide
(90% Removal)
179,000
2,542,000
262,000
848,000
3,831,000
520,000
$4,351,000
$2,412,000
$4,087,000
$50
2.30<7lb
Citrate
(95% Removal)
179,000
2,226,000
280,000
805,000
3,480,000
N/A
$3,490,000
$1,920,000
$3,268,000
$37
1.84e/lb
COMMENTS
Based on smelter rated capacity of 1100 TPD concentrate 25% copper, operating 340 days/year.
Annual production 89,000 TPY. Zero net-back to smelter for acid or sulfur produced.
-------
214
-------
12.5 KENNECOTT COPPER CORPORATION — GARFIELD, UTAH
This smelter is the largest of the Kennecott Copper Corporation
smelters with a nominal capacity of 2400 TPD of concentrate. Operations
began in 1907. Under the present operating configuration, green concen-
trate is fed to three reverberatory furnaces with the matte being
Processed through 6 or 7 on-line converters. Nine converters are avail-
at>le. Reverberatory gases are treated in the usual way through waste
boilers and an electrostatic precipitator before being vented out a
ft stack. The converter gases, after appropriate gas cleaning, feed
5 acid plants with a combined capacity of 1400 TPD acid.
As part of a compliance schedule submitted to the State of Utah,
Kennecott proposes to replace the reverberatory furnaces with 3 Noranda
factors with completion by early 1977. The approximately 10 percent
S°2 offgases will be directed to the acid plant system. The program
also includes the installation of water-cooled hoods on the 4 converters
required to balance the Noranda reactor system, and the provision of a
new 800 TPD acid plant to replace two of the present small, older units
and to increase rated acid production capability to 2000 TPD. A new
12°0 ft stack has been constructed as part of the Noranda reactor pro-
ject and is presently being lined although it will not be commissioned
the Noranda reactors are available.
Although the present reverberatory gas stream carries approximately
1 percent S02, no consideration will be given in this study to the
apPUcation of an absorption-based control system in view of the Noranda
teactor and related system changes and the July 1977 completion date.
215
-------
216
-------
!2.6 KENNECOTT COPPER CORPORATION ~ HAYDEN SMELTER, ARIZONA
12.6.1 Smelter Characteristics
The Kennecott Copper Corporation Hayden Smelter was put on stream
in 1958 with a single green-feed sidewall charged reverberatory furnace
and three Fierce-Smith converters. In 1969, a fluo-solids roaster was
conunissioned} with the cleaned effluent gases feeding a 750 TPD rated
sulfuric acid plant.
In January 1971, a comprehensive compliance plan was submitted to
the Arizona Air Pollution Hearing Board with accomplishment targeted for
early 1974. The program included improved reliability of the fluo-
s°lids reactor and gas handling system; installation of water-cooled
hoods and water spray coolers on each converter; new burner blocks,
Vater-cooled door frames and improved doors, and redesigned Wagstaff
feeders; elimination of the return of hot converter slag to the rever-
etatory furnace; and conversion of the 750 TPD acid plant to a double
c°ntact plant rated at 100,000 SCFM at 10.5 percent S02. Sections of
^e Program were completed and became operational during the plan period,
and the final acceptance tests on the acid plant were made in April
7^. in M^ 1974 , the Hayden Smelter of Kennecott Copper Corporation
ecame the first Arizona smelter to be issued a certified operating
Petmit, with total S02 control being better than the 90 percent reduction
by
Nominal capacity of the smelter is approximately 1100 tons/day of
Under the smelter's process configuration, the off gases from the
°aster, after mechanical cleaning in a cyclone system and cooling via
Wat-
6r sprays, pass through a venturi scrubber and a Peabody scrubber
ef°re mixing with the cleaned converter gases. The converter gases are
Ool*d by ultrasonic water sprays to 700°F and then pass through an
6ctrostatic precipitator and a Peabody scrubber. They are blended
the cleaned roaster gases, and the total stream passes through two
ailks of three parallel mist precipitators before being directed to the
^ le contact acid plant. Because of frequent periods of low gas
.
U8th or volume, a direct-fired preheater is kept on-line at all
s< Control dampers in the duct system vary the amount of process
217
-------
gas passing through the preheater, depending on the need of the plant.
The reverberatory gases, following usual practice, pass through a waste
heat boiler and electrostatic precipitator before being vented up a 600
ft stack.
The acid production is used within the Ray Mines Division of Kenne-
cott Copper,
Service facilities at this smelter—electric power and treated
water—are fully loaded. Water is available but is characteristically
hard.
12.6.2 Weak SO,, Streams
The flow schematic provided in Figure 12.6-1 indicates the gas flow
configurations for this smelter. The only weak S09 stream is that
produced by the reverberatory furnace. Only about 6-9 percent of the
total sulfur contained in the feed concentrate is released in the rever-
beratory and the concentration of SO™ in the resulting offgases is
accordingly very low. The gas stream is characterized in Table 12.6-1.
12.6.3 SOg Control Process Selection
Under existing regulations, further control of SO™ streams in this
smelter is not required and there is only an academic interest in se-
lecting an applicable SO™ control for the reverberatory stream.
As with most reverberatory gas streams, the relatively high oxygen
level of the Kennecott Hayden Smelter, together with the very low S02
concentration, poses a severe problem for most absorption-based control
systems. Oxidation of the usual sulfite or bisulfite absorption product
to sulfate is universally related to higher purge requirements, and
hence higher costs for makeup chemicals or for auxiliary recovery facilic
It is also frequently the cause for operational problems such as scaling*
Absorption processes which appear to be the least susceptible to these
oxidation effects are the citrate process and magnesium oxide process—
both regenerable type processes. However, since the latter process
produces a regeneration product which is a "dirty" and fairly weak (15
percent) S02 stream, and considering the small amount of S02 involved
(< 50 TPD), it does not appear to be a viable consideration for this
smelter.
The double alkali throwaway process may be also an applicable
process for this particular situation because of the limited quantity °*
S0_ involved and its ability to regenerate the sulfate routinely during
the active sodium ion regeneration sequence.
218
-------
SMELTER FLOW SCHEMATIC
1
r
CYCLONES
j FLUOSOLIDS
| ROASTER
I
I
I I
I I
COOLER
I I
I
I
to
— 1 1
1 1
1 1
1 1
SCRUBBER
75,000 SCFM (MAX)\
10% SO2 (MAX) J
\
i
1
1
1
I
i
1
I
I
I
i
! i
i
i
MAX 100,000 SCFM
MAX 10.5% SO2
I 1
TO ACID
PLANT
MIST
PRECIPITATORS
•CONVERTERSi
(3)
i
L,
(2 IN OPERATION
AT TIME)
I
I
'1
LI
COOLER
1 ESP f
1 J
I I
I I
L J
SCRUBBER
1150-165,000 SCFM
0.2-O.3% SO,
300°F
REVERB
(1)
WHB
ESP
^ STACK
"" 600 FT
FIGURE 12.6-1
-------
In considering the application of the citrate process, it is worth
noting that the Bureau of Mines is currently working on this process to
effect release of the SCL from the citrate absorbent using steam stripping*
Steam consumption is presently around 10 Ibs/lb SO- for 0.5 percent S0?
gas streams, but there is some expectation that this may be reduced. The
use of appropriate condensing and drying equipment would provide an
essentially 100 percent S02 stream which could be blended with a sulfuric
acid plant gas feed. This approach would be particularly effective for
the Kennecott Hayden Smelter. The existing plant with additional "standby'
capacity to accommodate the usual flow and S0~ fluctuations associated
with converter operation could handle the additional 300-400 SCFM of S02
via an appropriate control system responding to low S02 flow levels.
12.6.4 j>0_2 Control Process Costs
The citrate, modified citrate, and double alkali processes have
been selected as the potential control systems for this smelter and the
basic cost structures have been modified to accommodate the specific
conditions of this operation.
1) The temperature of the reverberatory gas stream at the
stack is 300°F, and the gas cooling and conditioning costs
have been reduced from these provided in Appendix A.
Capital costs for the identified maximum gas flow have been
reduced by the factor
T
o
R(300°F)
T°
0.6
R(600°F)
2) The retrofitting costs of making the necessary tie-ins of the
gas conditioning and S02 absorption columns with bypass and
dampers in the existing ductwork have been allowed for
by using a factor of 20 percent of the capital cost of the
gas conditioning and S02 absorption systems.
3) The variability of both gas flow and S02 content in the
reverberatory gases requires the provision of larger surge
capacity after the absorption system to provide a uniform
S02 rate to the regeneration or sulfur dioxide handling
section, and an additional 5 percent of the capital investfl»ett
of the S02 absorption section has been provided to cover
this modification.
220
-------
4) Additional service facilities, including an electrical
substation and a water treatment plant to support a
scrubbing process, have been included in the total
capital cost structure for each process. Costs have
been based on the power and water unit relationships
established for each process and the cost relationships
provided in Appendix F. A packaged boiler system and
treated water facility have been provided for the
citrate process S02 stripping option.
5) The citrate S02 stripping operation will also require
a suitably sized auxiliary S02 liquefaction and storage
plant to supplement the stripping facilities. Appendix
E provides capital and operating cost relationships
for the SCL liquefaction and storage plant.
6) To allow for those additional costs which are incurred
during any major construction program in an operating
plant, an additional 20 percent of the total boundary
limit cost for each process has been added to the capital
cost structure.
The capital and operating cost structures for the citrate process,
the citrate process plus its special S02 stripping option, and for
tlle double alkali process selected for the Hayden Smelter are provided
itl Tables 12.6-2 and 12.6-3, respectively. Individual computation
8heets for each process are provided under the appropriate smelter name
itl Appendix G.
221
-------
Table 12.6-1. KENNECOTT COPPER CORPORATION - HAYDEN SMELTER, ARIZONA.
CHARACTERIZATION OF WEAK S02 GAS STREAMS.
to
to
to
Stream
Reverb (1)
(at stack)
Volume
SCFM
150-165,000
Av. S02 Content
Sizing basis (Ibs/hr)
3000
Equivalent Av.
S02 Content (%)
0.2
% Oxygen (2)
10-15
Temp. °F
300
Notes. (1) No data on SO content available. Particulate
loading after ESP <0.02 grains/SCF.
(2) Oxygen content range estimated based on reported
test data of 13.0%.
-------
Table 12.6-2. KENNECOTT COPPER CORPORATION - HAYDEN, ARIZONA
CAPITAL COSTS FOR SCL CONTROL PROCESSES ON REVERBERATORY FURNACE OFF-GASES
ITEM
1. Boundary limits retro-
fitted primary system
2. Auxiliary Plant:
(a) S02 liquefaction
and storage
TOTAL Boundary Limit Cost
3. Total Service Costs
4. Associated Investment
Costs
(a) new ESP
(b) General site-
related costs
Total Capital Investment
Capital cost/annual ton
S0? Removed
Primary S02 Control Method
Double Alkali
(95% Removal)
6,120,000
N/A
6,120,000
675,000
1,224,000
8,020,000
$691
Citrate Process
(95% Removal)
7,660,000
N/A
7,660,000
820,000
1,530,000
10,010,000
$863
Citrate Process
with
Direct Stripping
5,890,000
200,000
6,090,000
930,000
1,220,000
8,240,000
$710
COMMENTS
-------
Table 12.6-3. KENNECOTT COPPER CORPORATION - HAYDEN, ARIZONA
ANNUAL OPERATING COSTS AND TOTAL NET ANNUALIZED COSTS
ro
-e-
ITEM
1. Gas conditioning and S02
absorption
2. SO- handling
3. Labor
4. Maintenance, Ins.& Taxes
5. Total Annual Operating
Cost
6. Auxiliary Plant Total
Operating Costs
(a) S0_ liquefaction
and storage
7. Total Annual Operating
Cost
8. Annuali zed Capital Cost
9. Total Net Annualized
Cost
10. Cost/year/ton SO-
Removed '
11. Cost/lb copper^ '
Primary S02 Control Method
Double Alkali
(95% Removal)
205,000
511,000
140,000
465,000
1,321,000
N/A
1,321,000
1,055,000
1,485,000
$128
l.OOc/lb
Citrate Process
(95% Removal)
223,000
240,000
245,000
588,000
1,296,000
N/A
1,296,000
1,316,000
1,670,000
$144
l.llC/lb
Citrate Process
with
Direct Stripping
223,000
441,000
175,000
481,000
1,320,000
15,000
1,335,000
1,084,000
1,514,000
$130
l.Oc/lb
COMMENTS
-------
12.7 KENNECOTT COPPER CORPORATION — HURLEY SMELTER, NEW MEXICO
12.7.1 Smelter Characteristics
This smelter, which began operations in 1939, is a relatively small
smelter with a nominal capacity of 850 TPD of reverberatory feed. Green
concentrate is presently processed in one of two available reverberatory
furnaces with the matte being processed through three converters.
Offgases from the reverberatory pass through a waste heat boiler, are
treated in an electrostatic precipitator, and are then vented to atmosphere
through a 500 ft stack. The reverberatory and flue system is in very
poor condition, and air in-leakage is excessive. The electrostatic
precipitator cannot adequately handle the accompanying gas flow rates
with the result that efficiency levels are low, estimated at under 60
percent. Maximum average grain loading has been reported as 0.97 gr/SCF.
Offgases from the converters (3 normally in operation) were originally
cleaned of particulate material in a battery of multiclones and vented
up a 625 ft stack. In December, 1974, a converter gas collection system
with water cooled hoods and a 650 TPD rated double contact sulfuric acid
plant with the usual gas cleaning facilities was commissioned to process
these converter gases. The acid plant has a maximum flow capability of
around 90,000 SCFM at approximately 4.5 percent S02 content.
Kennecott in a variance petition dated March 1, 1974, has requested
relief through October 1, 1978, for its reverberatory stack particulate
emissions on the basis of their research and development program to
implement an improved smelting process based on converter smelting with
oxygen-enriched blast air and supplemental heating. The resulting white
metal would then be transferred to a finishing converter for production
of blister copper in the usual manner. If successful, this converter
smelting process will essentially eliminate the reverberatory furnace
process. A report on the status of this program and Kennecott's future
action is due to the State of New Mexico Environmental Improvement
Agency in August, 1975.
The smelter is located about 7 miles from the mine site, but land
availability at the smelter is not a problem and a one square mile
tailing pond is available. The smelter has its own supply of limestone
and an on-site lime plant although its capacity is limited to the
present needs of the smelter and reduction plant.
225
-------
A source of good quality water is available but additional pumping
facilities would be needed to make it available at the smelter site. The
existing power system appears to be capable of meeting the additional
demand from an SCL control system.
12.7.2 Weak S02 Streams
The weak SO^ stream in the Hurley smelter results from the rever-
beratory furnace operation and the high degree of dilution which occurs
down the gas collection system. A flow schematic is provided in Figure
12.7-1. With a nominally uniform green reverberatory charge, furnace
operating conditions are relatively stable. The charging operations and
"soot" blowing of the waste heat boilers contribute to peak levels in
both gas volume and S02 content but the excessive dilution taking place
down the collection system tends to minimize these disturbances. The
gas stream has been characterized at the stack breeching as shown in
Table 12.7-1.
12.7.3 S(X, Control Process Selection
Although this smelter is doing research and development to eliminate
its weak S02 stream, for the purposes of this study, the present rever-
beratory gas stream as characterized in Table 12.7-3 will be considered
as the feed to an absorption-based SO- control process.
The high degree of dilution with the resulting low level of S02 a°d
high oxygen content of the final effluent reverberatory gas stream
generally makes such a stream unfavorable for the application of liquid
scrubbing control systems. The higher expected oxidation rates suggest
that the Wellman-Lord, ammonia, xylidine, and even the limestone process
would all be penalized significantly by both higher operating costs
through increased purge requirements and more difficult operating con-
ditions. The direct application of a sulfuric acid plant with preheati0*
and refrigeration is obviously an unfavorable approach. The citrate a°
magnesium oxide processes are less sensitive in the final analysis to
oxidation and are applicable processes, although the magnesium oxide
process may be susceptible to higher scaling in the absorber and may
also require magnesium sulfate purge and recovery facilities. Energy
costs and the additional requirement for gas cleaning equipment and a°
226
-------
KENNECOTT COPPER — HURLEY, NEW MEXICO
SMELTER FLOW SCHEMATIC
REVERB
(2)
WHB
240-270,000 SCFM
0.6-0.82% S02
17% O2
425° F
ESP
(1 IN OPERATION
AT TIME)
N>
Ni
STACK
500 FT
I 1
, CONVERTER }_
{ (4) '
(3 IN OPERATION
AT TIME)
I I
I ACID PLANT I
1 (INCLUDING GAS | STACK
I
I
I
L_
CONDITIONING) |
I
I
750 TPD
FIGURE 12.7-1
-------
auxiliary sulfuric acid plant associated with this latter process are
additional adverse characteristics. The double alkali system will incur
higher oxidation through the high oxygen levels of the reverberatory
stream but the lime treatment in the regeneration step of the dilute
system converts the Na^SO, product to sodium hydroxide and calcium
sulfate, with the calcium sulfate being precipitated along with the
other regeneration precipitates.
The cost structures for the double alkali, citrate and magnesium
oxide processes selected for the Kennecott, Hurley, New Mexico smelter
including support services and miscellaneous charges are provided in
Tables 12.7-2 and 12.7-3.
12.7.4 S02 Control Process Costs
The basic cost structures for the selected processes for this
smelter have been modified to take into consideration those special
factors and conditions associated with this operation.
1) The temperature of the reverberatory gas stream at the stack
is 430-450°F and the gas cooling and conditioning requirements
will be reduced from those specified in Appendix A. Capital
costs for the identified maximum gas flow have been reduced
by the factor
0.6
R(450°F)
T R(600°F)
2) The retrofitting costs of making the necessary tie-ins of the
gas conditioning and SO™ absorption columns with bypass and
dampers in the existing ductwork have been allowed for by
using a factor of 25 percent of the capital cost of the gas
conditioning and SCL absorption systems.
3) The existing electrostatic precipitator at Hurley is not
adequate for the gas flows it is presently handling and the
installation of an absorption based SCL control process would
necessitate the installation of a replacement ESP sized to
handle approximately 500,000 ACFM at 500°F. However, since
the existing unit is not presently meeting the particulate
emissions regulation, its replacement is not dependent on
the installation of the absorption system and the cost is
not chargeable against the SO, control system itself.
228
-------
4) An additional 5 percent of the capital investment of the
S0~ absorption section has been provided to allow for
the larger surge capacity after the absorption section to
accommodate the variable SO™ content and gas flow rates
and to provide a uniform SO- rate to the regeneration section.
5) An allowance has been provided for a new water pumping station
but other existing service facilities appear to be capable
of handling the requirements of a scrubbing system.
6) A general site cost factor of 20 percent of the total
boundary costs for each process has been allowed to cover the
additional cost associated with the implementation of major
process changes in an operating plant.
Individual calculation sheets for the applied processes are provided
under the appropriate smelter heading in Appendix G. A summary of these
capital costs and associated annual operating and total annualized costs
is provided in Tables 12.7-2 and 12.7-3, respectively.
229
-------
Table 12.7-1. KENNECOTT COPPER CORPORATION - HURLEY SMELTER, NEW MEXICO
CHARACTERIZATION OF WEAK SO GAS STREAMS.
Stream
Reverb ^
Volume
SCFM
240-270,000
Av. S02 Content'1'
Sizing basis (Ibs/hr)
20,200
Equivalent Av.
S02 Content (%)
0.74
% Oxygen(2)
15-18
Temp. °F
425
CO
O
Notes. (1) No data on particulate loading available
SO content 0.05%.
(2) Oxygen content of gas stream estimated.
-------
Table 12. 7-2. KENNECOTT COPPER CORPORATION - HURLEY SMELTER, NEW MEXICO
CAPITAL COSTS FOR SO2 CONTROL PROCESSES ON REVERBERATORS FURNACE OFF-GASES
ITEM
1. Boundary limits retrofitted
primary system
2. Auxiliary Plant:
(a) H SO, plant with dry
gas cleaning
TOTAL Boundary Limit Cost
3. Total Service Costs
4. Associated Investment Costs
(a) new ESP
(b) General site-related
costs
Total Capital Investment
Capital cost/annual ton S0~
Removed
Primary S0_ Control Method
Double Alkali
(95% Removal)
12,080,000
N/A
12,080,000
50,000
( 3,500,000)
2,420,000
$14,550,000
$186
Magnesium Oxide
(90% Removal)
14,150,000
N/A
14,150,000
50,000
( 4,000,000)
2,830,000
$17,030,000
$217
Citrate
(95% Removal)
12,900,000
3,900,000
16,800,000
50,000
( 4,000,000)
3,360,000
$20,210,000
$266
COMMENTS
Gas cleaning by high temperature
bag filters. S09 removal effici-
ency 97%
Additional pumping facilities
only.
This item not chargeable to
S0£ control.
N>
CO
-------
Table 12.7-3. KENNECOTT COPPER CORPORATION - HURLEY SMELTER, NEW MEXICO
ANNUAL OPERATING COSTS AND TOTAL NET ANNUALIZED COSTS
ITEM
1. Gas conditioning and S0_
absorption
2. SO- handling
3. Labor
4. Maintenance, Ins. & Taxes
5. Total Annual Operating Cost
6. Auxiliary Plant Total
Operating Costs
(a) Sulfuric acid plant
7. Total Annual Operating Cost
8. Annualized Capital Cost
9. Total Net Annualized Cost
10. Cost/year/ton SO^ removed
11. Cost/lb copper^3'
Primary S0_ Control Method
Double Alkali
(95% Removal)
393,000
3,439,000
210,000
888,000
4,930,000
N/A
4,930,000
1,937,000
4,029,000
$51
2.42c/lb
Magnesium Oxide
(90% Removal^
440,000
2,169,000
315,000
1,065,000
3,989,000
470,000(2)
4,459,000
2,683,000
4,104,000
$57
2.47c/lb
Citrate
(95% Removal)
440,000
1,601,000
1,047,000
3,421,000
N/A
3,421,000
2,263,000
3,491,000
$45
2.10c/lb
COMMENTS
to
UJ
10
(1)
(2)
Overall S02 removal with H SO plant is 87.3%.
Includes maintenance, insurance, and taxes.
/0\
Based on smelter rated capacity of 850 TPD concentrate. Reverberatory fuel includes 15% precipitate
copper to give total copper content of charge 30.3%, 340 days/year operation. Annual production
83,000 TPY. Zero net-back to smelter for acid or sulfur produced.
-------
12.8 KENNECOTT COPPER CORPORATION -- McGILL SMELTER, NEVADA
12.8.1 Smelter Characteristics
This smelter began operations in 1907. Two reverberatory furnaces
currently handle approximately 800-1000 tons/day of copper concentrate
with the output being processed through three converters. Offgases from
the reverberatories pass through waste heat boilers and electrostatic
Precipitators before being vented out a new 750 ft stack. The electrostatic
Precipitators are old and recovery efficiencies are low (70-80 percent).
°ffgases from the converters are treated in an electrostatic precipitator
and are presently vented up the same 750 ft stack as the reverberatory
gases.
A compliance schedule with completion originally scheduled for mid-
has been submitted encompassing upgrading of the reverberatory
's to a proposed 98.9 percent, installation of water-cooled hoods on
the converters and the provision of a 500 TPD single contact sulfuric
acid plant. Tail gas would be vented up a separate 250 ft stack with
the 750 ft stack handling the reverberatory gases. Emissions collected
fr°m auxiliary hooding in the converter area would be vented to the
reverberatory furnace gas system prior to the stack. At the same time
CaPacity of the smelter would be reduced from the nominal capacity of
1000 TPD of concentrate to approximately 750 TPD by the elimination of
custom smelting business. Implementation of this plan is presently
abeyance pending an appeal by Kennecott for Intermittent Control
(ICS) options. The State of Nevada's S02 emission standard calls
f°r 60 percent control of the input sulfur, and approximately 64 percent
Cotitrol would be accomplished via the proposed changes. However, as
n°ted by the State authorities, existing ambient air standards are not
aPParently being violated with the present level of S02 emissions.
This smelter is located in a rather harsh environment at an elevation
of 600-7000 ft. Water, which is in short supply, is piped from a source
5 miles away. This water shortage has caused problems in the operation
°f their tailings pond. Although the size of the pond has been reduced
to approximately 2 square miles as part of the effort to control dust,
e original size of the pond was approximately 6 square miles.
233
-------
12.8.2 Weak SO., Streams
Under the present operating mode, both the reverberatory and converter
offgases are weak SQ~ streams. A flow schematic is provided in Figure
12.8-1 based on the present operating level of approximately 1000-1100
TPD concentrate, with the converter gases being vented to the stack.
For the purposes of this study, it is assumed that the proposed
acid plant and the upgraded converter gas system will be installed as
provided for in the compliance schedule. However, the proposed reduction
of the smelter capacity will not be taken into consideration, and the
gas streams will be characterized on the basis of present capacity gas
flow levels as provided in Table 12.8-1.
12.8.3. S00 Control Process Selection
J £, """'" "" ' ^ L »....„
The elevation of this smelter, together with the high oxygen levels
in the reverberatory gases, are not favorable conditions for an absorp-
tion based SCL control process. The effect of the 6000-7000 ft ele-
vation will be to reduce absorption efficiency, and to compensate,
additional cooling of the absorbent liquid, higher liquid/gas (L/G)
ratios, or additional absorption stages may be necessary. The higher
oxygen levels in the gas stream will produce higher oxidation levels in
the absorbent with significantly higher purge requirements and loss of
active chemicals, or alternatively, greater capital investment for purge
treatment and recovery facilities.
Under these conditions only the double alkali, citrate and the
magnesium oxide processes appear to offer reasonable potential as appli"
cable control systems.
12.8.4 S09 Control Process Costs
'"""""'/, ' " ' " ' ' •••'---•"--"-
The absorption characteristics of the double alkali, citrate and
the magnesium oxide processes are such that the high elevation of this
smelter will have only a moderately negative effect; and no special cost
allowance to compensate for lower absorption efficiency will be provided
The following conditions specific to this smelter will impact on the
capital and operating cost structure.
234
-------
KENNECOTT COPPER — McGILL, NEVADA
SMELTER FLOW SCHEMATIC
200-250,000 SCFM
0.94% SO2
500° F
REVERB
(2)
WHB
ESP
fc-
1
STACK
750 FT
CONVERTERS
•» I" WATER COOLED"]
-I ESP j
I I
HOODS PROPOSED J
I H2S04 PLANT
T 500 TPD I
] (S.C.)
I 1
PROPOSED
STACK
250 FT
FIGURE 12.8-1
-------
1) Temperature of the reverberatory gases after the
electrostatic precipitators is between 400-500°F,
and gas cooling and conditioning requirements would
be reduced from the 600°F temperature basis selected
for the general cost relationship provided in Appendix
A. Using an average temperature of 450°F, costs will
be adjusted on the basis of
T°R(450°F)
T°R(600°F)
0.6
2) The additional costs incurred by installing and tieing
in the gas conditioning and SO™ absorption sections to
existing flues and providing bypasses and appropriate
dampers have been allowed for by the use of a retrofit
factor of 30 percent of the capital investment of these
sections.
3) To accommodate the variability in both gas flow and SO^
content in the reverberatory gases, larger surge
capacity after the absorption system will be necessary
to provide a uniform S0« rate to the regeneration section.
An additional 5 percent of the capital investment of
the SO- absorption section has been provided.
4) Additional service facilities including substations, power
distribution system, and water treatment plant to support a
scrubbing process will be necessary; and estimated costs
have been based on the unit values established for each
control process and the cost data provided in Appendix F.
It should be noted that the cost of a cooling tower has
been included in the costs developed for the gas
conditioning section.
5) An auxiliary sulfuric acid plant taking an 8 percent S02
stream feed has been provided with the magnesium oxide
process. Capital and operating costs of this plant based
on the S02 hourly rate of the magnesium oxide process itself
are provided in Appendix C.
6) As with all on-going operating plants, major process changes
incur substantial additional costs related to coordination
and scheduling difficulties, space limiations, special
safety requirements, etc. To allow for these not readily
identifiable costs, an estimate of 20 percent of the total
boundary limit costs for each process has been allowed.
7) The cost of replacing the present electrostatic precipitate*
is not allocated against the S02 control system costs.
Replacement has already been planned.
236
-------
The capital and operating cost structures for the three processes
identified for the McGill Smelter—double alkali, magnesium oxide,
and citrate—are provided in Tables 12.8-2 and 12.8-3, respectively.
Individual computation sheets for each process are provided under the
appropriate smelter name in Appendix G.
237
-------
Table-12,8-1. KENNECOTT COPPER CORPORATION - McGILL SMELTER, NEVADA.
CHARACTERIZATION OF WEAK S02 GAS STREAMS.
Stream
Reverbs. (2)
Volume
SCFM
200-250,000
(210,000 Av.)
Av. SO Content
Sizing basis (Ibs/hr)
20,000
Equivalent Av.
S02 Content (%)
0.94
% Oxygen (2)
12-16
Temp. °F(3)
500
to
u>
oo
Notes. (1) No data on S03 loading but average participate
loading 0.35 gr/cubic foot.
(2) Oxygen content of gas stream estimated.
(3) Temperature at stack itself 400°F.
-------
Table 12.8-2. KENNECOTT COPPER CORPORATION - MCGILL, NEVADA
CAPITAL COSTS FOR SO2 CONTROL PROCESSES ON REVERBERATORY FURNACE OFF-GASES
ITEM
1. Boundary limits retrofitted
primary system
2. Auxiliary Plant:
(a) H2S04 plant with dry
gas cleaning
TOTAL Boundary Limit Cost
3. Total Service Costs
4. Associated Investment Costs
(a) new ESP
(b) General site-related
costs
Total Capital Investment
Capital cost/annual ton
SO- Removed
Primary S02 Control Method
Double Alkali
(95% Removal)
12,250,000
N/A
12,250,000
780,000
2,450,000
$15,480,000
$200
Magnesium Oxide
(90% Removal)
13,010,000
3,950,000
16,960,000
1,000,000
3,390,000
$21,350,000
$300
Citrate
(95% Removal)
14,290,000
N/A
14,290,000
980,000
2,820,000
$18,090,000
$233
COMMENTS
to
CO
-------
Table 12.8-3. KENNECOTT COPPER CORPORATION - MCGILL, NEVADA
ANNUAL OPERATING COSTS AND TOTAL NET ANNUALIZED COSTS
/
ITEM
1. Gas conditioning and SO-
absorption
2. S02 handling
3 . Labor
4. Maintenance, Ins. & Taxes
5. Total Annual Operating Cost
6. Auxiliary Plant Total
Operating Costs
(a) Sulfuric acid plant
7. Total Annual Operating Cost
8. Annualized Capital Cost
9. Total Net Annualized Cost
10. Cost/year/ton S0_ Removed
11. Cost/lb copper ^
Primary SO- Control Method
Double Alkali
(95% Removal)
356,000
3,404,000
210,000
908,000
4,878,000
N/A
$ 4,878,000
$ 2,036,000
$ 4,077,000
$53
3.16c/lb
Magnesium Oxide
(90% Removal)
400,000
2,210,000
315,000
1,098,000
4,023,000
470,000
$ 4,493,000
$ 2,807,000
$ 4,461,000
$61
3.45c/lb
Citrate
(95% Removal)
'400,000
2,035,000
333,000
1,069,000
3,837,000
N/A
$ 3,837,000
$ 2,379,000
$ 3,795,000
$49
2.94c/lb
COMMENTS
to
*>
O
Based on smelter operating capacity of 800 TPD concentrate. Reverberatory feed includes
8 % precipitate copper to give total copper content of charge = 25%. 340 days/year
operation. Zero net-back to smelter for acid or sulfur produced.
-------
12.9 MAGMA COPPER COMPANY — SAN MANUEL SMELTER, ARIZONA
12.9.1 Smelter Characteristics
The Magma Copper Smelter began operation in 1956. The smelter is a
relatively modern facility with a present capacity of approximately 2500
TPD concentrate. Completion of long range plans will ultimately raise
the capacity of this smelter to approximately 2900 TPD concentrate. The
smelter has made substantial investments in pollution control equipment
and modifications over the past several years. Green concentrate is
smelted in three reverberatory furnaces, with the offgases being treated
in new high efficiency electrostatic precipitators before venting out a
515 ft stack. Final completion of the installation of these ESP's is
scheduled for late 1975. Output from the reverberatory furnaces is
treated in 5 ,Pierce-Smith converters. A sixth converter is presently
eing installed with completion expected at the end of 1975. New
ucting and water-cooled hoods were installed on the converters over
73-1974, and the entire gas flow now goes to a single contact, 2-train
Sulfuric acid plant rated at 2480 TPD. The parallel trains provide
UlUisual flexibility. Although the 93 percent acid is used internally in
toall quantities with a significant amount being marketed, a supple-
ental acid neutralization plant using basic tailings and limestone has
een provided. This facility can accept the total output of the acid
ant for short periods of up to 2-3 days if necessary. Approximately
tons/day of acid are being neutralized at the present time due to
Cess supply over demand. Although actual data are not available on
e Performance of the new ESP's currently being installed on the rever-
ratory furnace offgas line, it is expected that particulate loading
H be reduced to below 0.02 grains/SCF.
The smelter utilizes waste heat boilers in the reverberatory offgas
^es to generate in-house power, with supplemental purchases from
lie service facilities transmitting over long distances. The extreme
aness and high salt content of available water require extensive
tteatment.
12 9 0
Weak SOU Streams
The weak S00 stream in this smelter is produced by the reverberatory
^tn
aces. A flow schematic is provided in Figure 12.9-1. The reverberatory
flv
§e is uniform and only minor variations in the fuel/charge ratio
241
-------
As with all reverberatory offgases, the operations of charging and
"soot" blowing of the waste heat boilers, together with normal air
infiltration along the gas handling system, have a significant effect on
both gas volumes and SCL concentrations. Total air infiltration may
range as high as 100-200 percent over a typical operating cycle which
includes charging the furnace, smelting the concentrate, and "soot"
blowing the waste heat boilers. From the data available, and based on
the present operating capacity, the reverberatory gas stream at the
stack breeching has been characterized at the representative values
indicated in Table 12.9-1.
12.9.3 S00 Control Process Selection
11 ' ' Z, • "' . _. --—--.._.
The application of an SO,, control process based on gas scrubbing to tb*
reverberatory gases of the Magma Copper Company faces the same uncertainti6
noted previously—wide fluctuations in both gas volumes and SC^ contents,
and high oxygen contents. The gas handling sections—gas conditioning
and SCv absorption—must be sized for the maximum expected gas flow
independent of the varying range of S02 in the gas stream. The S02
handling section of the process must be sized to handle an "average" S02
rate established via appropriate surge capacity after the S02 absorption
section.
The high oxygen levels associated with these reverberatory gas
streams limit the S09 control processes which might be considered as
appropriate matches for this smelter. High oxygen levels will generally
result in increased oxidation rates in the absorbent thereby producing
sulfates which must be purged from the system. High purge rates and the
associated loss of the active absorbent may inflate the annual operating
costs to unrealistic levels, or alternatively, increase capital invest-
ment and operating costs by necessitating the provision of secondary
recovery and treatment facilities.
The lime/limestone throwaway system does not appear to be applicabl6
under these conditions. The Smelter Control Research Association (SCRA)
sponsored limestone scrubbing pilot plant which operated in 1972 on a
1.0 percent SO- content reverberatory gas level was plagued with severe
scaling and operational problems. The variable S02 level and the 10-15
percent oxygen level in the Magma Smelter gases suggest that similar
242
-------
MAGMA COPPER COMPANY — SAN MANUEL, ARIZONA
SMELTER FLOW SCHEMATIC
400-550,000 SCFM
0.7% AV S02
10-15% 02
400-550°F
REVERBS
(3)
WHB
ESP
STACK
515 FT
ISi
J>
co
NEW BEING I
[INSTALLED!
I 125-175,000 SCFM
40-5.5% SOo
I 1
'CONVERTERS J
r
i
I —
I
-I ESP \
I !
I 1
GAS !
COND. |"
I 1
STACK
2500 TPD ACID PLANT S.C.
*• STACK
(2 PARALLEL SECTIONS)
FIGURE 12.9-1
-------
problems might be expected, and the process cannot be considered a
suitable application. Although the double alkali process is also sus-
ceptible to the high oxygen content and related oxidation, its re-
generation chemistry in the dilute absorbent mode provides sulfate
regeneration and the process appears to offer an acceptable control
approach for this smelter. However, it is noted that a new SCRA pilot
plant project using the ammonia double alkali system started up at San
Manuel in May 1975, and this program may yield more detailed information
on oxidation characteristics.
Of the regenerable processes, only the magnesium oxide and citrate
processes appear to have the capability of treating high oxygen content
SCL streams without substantial penalty. In the magnesium oxide process»
sulfate formation may be self-limiting at around 15 percent MgSO,;
otherwise, purging and secondary treatment of the MgSO, will be necessary*
If the MgSO, concentration increases to 33 percent weight percent,
MgSO,-7H20 is precipitated with the MgSO *6H20 and the sulfate can be
reduced in the calcination step using coke. The high energy require-
ments of this process and the production of a relatively weak (15 per-
cent) S02 stream which needs cleaning and conditioning prior to an
auxiliary sulfuric acid or elemental sulfur plant are disadvantages.
The citrate process is relatively insensitive to oxidation, while its
absorption characteristics are such that fluctuating S02 levels can be
handled readily and effectively. Although this process is still under
development, it appears to offer the most potential of the candidate
processes for this smelter. In 1972, a small pilot plant using the
citrate process was operated at San Manuel, but gave discouraging
results in terms of citrate losses. Under subsequent development work
carried out by the Bureau of Mines, this deficiency appears to have bee*1
largely overcome.
There is also a possibility of using the citrate process with S*^
regeneration only via steam stripping and using the S02 from liquefied
storage to increase the S02 concentration in the converter gas
However, steam consumption values appear to be on the order of 8
S02 for this level of S02 removal, and this must be considered a maj°r
disadvantage. The Bureau of Mines is currently working on this aspect
of the process in an attempt to reduce steam requirements.
244
-------
Considering the gas volume and the S02 concentrations involved, the
direct application of a sulfuric acid plant to the Magma Copper Smelter
reverberatory gases is not a realistic alternative.
12.9.4 SOg Control Process Costs
Special factors or conditions associated with the Magma Smelter
which modify the basic cost structures of the selected processes are as
follows:
1) The temperature of the reverberatory gas stream at the stack
is 400-550°F. 500°F has been taken as an average value,
and the capital costs from Appendix A for gas cooling and
conditioning the identified maximum gas flow have been
modified by the factor
T°R(500°F)
T
o
R(600°F)
2) At Magma, the reverberatory gas stack is located very close
to the electrostatic precipitator itself, and a flue length
is only approximately 25 feet. Auxiliary equipment located in
this area would have to be relocated. Effecting the necessary
tie-ins of the gas conditioning and SO- absorption equipment
in this situation would be moderately difficult, and a factor
of 25 percent of the capital cost of these two sections has
been allowed to cover the additional cost.
3) The variability of both gas flow and SO, content in the
reverberatory gas stream will require the provision of larger
surge capacity after the absorption system to provide a
uniform S0? rate to the regeneration section, and an additional
5 percent of the capital investment of the SO^ absorption
section has been provided to cover this modification.
4) Additional electrical substation and distribution facilities
and a water treatment plant to support the makeup water
requirements of the scrubbing system have been provided, with
the estimated cost being based on the unit relationships
established for each process and the cost curves in Appendix
F.
5) The implementation of a major process change in an operating
plant incurs significant additional costs related to
scheduling difficulties, special safety and operating require-
ments, etc. An estimate of 15 percent of the total boundary
limit cost for each process has been allowed.
245
-------
The capital and operating cost structures for the three processes
selected for the Magma Copper Smelter—double alkali, magnesium oxide
and citrate processes—are provided in Tables 12.9-2 and 12.9-3,
respectively. Individual computation sheets for each process are
provided under the Magma Copper name in Appendix G.
246
-------
Table 12.9-1. MAGMA COPPER COMPANY - SAN MANUEL SMELTER, ARIZONA
CHARACTERIZATION OF WEAK SO GAS STREAMS.
Stream
Reverbs. (3)
Volume
SCFM
400-550,000
Av. S02 Content
Sizing basis (Ibs/hr)
34,500
Equivalent Av.
S02 Content (%)
0.7
% Oxygen(2)
10-15
Temp. °F
400-550
10
OS
Notes. (1) No data on 803 available. Particulate
loading less than 0.02 grains/SCF.
(2) Oxygen content of gas stream estimated.
-------
Table 12.9-2. MAGMA COPPER COMPANY - SAN MANUEL, ARIZONA
CAPITAL COSTS FOR S02 CONTROL PROCESSES ON REVERBERATORY FURNACE OFF-GASES
ITEM
1. Boundary limits retrofitted
primary system
2. Auxiliary Plant:
(a) H2SO^ plant with dry
gas cleaning
TOTAL Boundary Limit Cost
3. Total Service Costs
4. Associated Investment Costs
(a) new ESP
(b) General site-related
Total Capital Investment
Capital cost/annual tons SO^
Removed
Primary SO- Control Method
Double Alkali
(95% Removal)
17,980,000
N/A
17,980,000
1,160,000
2,700,000
$21,840,00
$163
Magnesium Oxide
(90% Removal)
19,810,000
5,650,000
25,460,000
1,230,000
**•
3,820,000
$30,510,000
$228
Citrate
(95% Removal)
21,060,000
N/A
21,060,000
1,330,000
3,160,000
$25,550,000
$191
COMMENTS
ro
*»
oo
-------
Table 12.9-3. MAGMA COPPER COMPANY - SAN MANUEL, ARIZONA
ANNUAL OPERATING COSTS AND TOTAL NET ANNUALIZED COSTS
ITEM
1. Gas conditioning and S0_
absorption
2. SO handling
3. Labor
4. Maintenance, Ins. & Taxes
5. Total Annual Operating Cost
6. Auxiliary Plant Total
Operating Costs
(a) Sulfuric acid plant
7. Total Annual Operating Cost
8. Annualized Capital Cost
9. Total Net Annualized Cost
10. Cost/year/ton S0_ Removed
11. Cost/lb copper' '
Primary S02 Control Method
Double Alkali
(95% Removal)
833,000
5,876,000
210,000
1,315,000
8,234,000
N/A
$ 8,234,000
$ 2,872,000
$ 6,455,000
$48
1.990/lb
Magnesium Oxide
(90% Removal)
942,000
3,835,000
315,000
1,612,000
6,704,000
690,000
$ 7,394,000
$ 4,012,000
$ 6,881,000
$56
2.12c/lb
Citrate
(95% Removal)
942,000
2,722,000
333,000
1,556,000
5,553,000
N/A
$ 5,553,000
$3,360,000
$ 5,430,000
$41
1.68c/lb
COMMENTS
ts»
js
vO
(1)
Based on smelter rated capacity of 2,500 TPD concentrate, 20% copper operating 340 days/year
Annual production 162,000 TPY. Zero net-back to smelter for acid or sulfur produced.
-------
250
-------
12.10 PHELPS DODGE CORPORATION — AJO, ARIZONA
12.10.1 Smelter Characteristics
The Ajo plant is a 25 year old copper smelter, handling about 700
TPD of concentrate. Uniform green charge is fed to a single rever-
beratory furnace. Matte is processed in three converters. The rever-
beratory offgases are processed through a waste heat boiler and an
electrostatic precipitator which removes particulates to a level of 0.06
gr/ACF. The converter gases are similarly cleaned in an ESP before
entering the gas conditioning section of a single absorption acid plant
rated at 750 tons of sulfuric acid per day. The tail gas from the acid
Plant is vented to a stack. The mining and concentrating plants adjoin
the smelter, which makes possible conveyor feed to the reverberatory.
Belt slingers feed the furnace from the sides.
12.10.2 Weak SO,. Stream
The only weak stream originates from the reverberatory furnace. A
flow schematic is provided in Figure 12.10-1. This gas has been vented
to the atmosphere but it is planned to treat this stream in a dimethy-
^aniline (DMA) scrubbing system which is expected to start up around the
Diddle of 1975. The gas stream has been characterized in Table 12.10-1.
12'10.3 S00 Control Process Selection
£,
An SO, control system for the reverberatory gases has already been
Elected for the Ajo smelter by the installation of a DMA scrubbing
Astern completed in 1972. Originally, it was to handle a mixture of
reverberatory and converter gases but equipment and operating problems
Prevented effective application. The plant has since undergone consider-
able development and modification. It is now planned to treat the
reverberatory gases only and to utilize less than the designed capacity.
Startup of the plant under these conditions is expected in mid-1975.
Although this process has been previously applied to copper smelter
converter gases in the 4-10 percent S02 range, the Ajo installation will
6 Dandling a much weaker gas stream with S02 concentrations in the 1-3
Percent range. There is some uncertainty as to the operating effectiveness
°f the DMA Process at these low S09 levels.
251
-------
12.10.4 S02 Control Process Costs
The original design basis for this installation was about 168 TPD
of liquid SO* but it will operate considerably below this capacity when
it comes into operation on the reverberatory gases. At its original
design base it was similar in capacity to ASARCO's 200 TPD rated plant
in the Tacoma, Washington smelter; but the reported costs for this
latter unit excluding gas conditioning appear to be substantially greater
than the Ajo installation. Although the plant has undergone changes and
modifications since 1972, including the addition of electrostatic mist
precipitators following the original humidifying and cooling towers, the
total cost including gas conditioning has been reported as $5,400,000.
If this cost is assumed to be averaged at 1971 cost levels, the escalated
value at mid-1974 is approximately $7,000,000. From the capital cost
curves presented in Figure 8-2 (based largely on ASARCO reported costs)
and including an adjusted gas conditioning cost for the Ajo temperature
of 450°F, the total capital cost would be approximately $8,800,000
including a retrofitting allowance. Because the costs have been based
on an actual smelter installation, an additional site cost allowance has
not been provided as with the other control processes.
Operating costs have been calculated from the unit usage and cost
data presented in Table 8.1. It must be noted however, that these costs
are oriented towards the xylidine system, and accordingly, they report
higher usage values than may be experienced with a dimethylaniline
system.
Table 12.10-2 provides a summary of both capital and annual costs
based on the cost relationships developed for the DMA/xylidine process
in Section 8. The computations for the annual operating costs are
included in Appendix G.
252
-------
PHELPS DODGE - AJO, ARIZONA
SMELTER FLOW SCHEMATIC
to
Ol
u>
I CONVERTERS l-
(3)
45-55,000 SCFM
1-3% SO2
REVERB
(1)
WHB
ESP
_! GAS !
! CLEANING I
DMA
rLANT
*
60-78.000 SCFM
4-9% SO2
I T
I i
-J ESP
...
—I GAS |_
J CLEANING I
ACID
PLANT
750 TPD
STACK
TO LIQUID
SO2 STORAGE
STACK
FIGURE 12.10-1
-------
Table 12.10-1. PHELPS DODGE CORPORATION - AJO SMELTER - ARIZONA
CHARACTERIZATION OF WEAK SOO GAS STREAMS
Stream
Reverb .
Volume
SCFM
45-55,000
Av. S02 Content
Sizing basis (Ibs/hr)
10,100
Equivalent Av.
SO™ Content (%)
2.0%
(3)
% Oxygen Feed
to DMA Plant
8 - 10%
Temp. °F
450°
N>
NOTES: (1) Particulate loading after ESP has been indicated
as 0.02 grains/AFCM .
(2) SO, content at the furnace uptake has been
indicated as 0,1%.
(3) Oxygen content of gas stream at stack breeching estimated.
-------
Table 12.10-2. PHELPS DODGE CORPORATION - AJO, ARIZONA
CAPITAL AND ANNUAL OPERATING COSTS FOR DMA/XYLIDINE
CONTROL SYSTEM ON REVERBERATORY FURNACE OFF-GASES
TOTAL CAPITAL INVESTMENT
Capital Cost/Annual tons
SO ~ Removed
ANNUAL DIRECT OPERATING COST
Gas Conditioning
SO- Absorption & Handling
Labor
Maintenance, Ins. & Taxes
ANNUALIZED CAPITAL COST
TOTAL NET ANNUALIZED COST
ANNUAL COST/TON so2 REMOVED
COST/LB COPPER
$ 62,000
832,000
140,000
589,000
$8,830,000
$223
$1,623,000
$1,161,000
$1,722,000
$ A3
1.57e/lb
Based on smelter rated capacity 700 TPD concentrate, 24% copper
operating 340 days/year. Annual production 55,000 TPY. Zero net-back
to smelter for SO. or acid produced.
255
-------
256
-------
12.11 PHELPS DODGE CORPORATION — DOUGLAS, ARIZONA
12.11.1 Smelter Characteristics
The Douglas Plant dates back to around 1903. It has a rated green
concentrate capacity of about 2800 TPD although the present operating
level is around 2000 TPD concentrate. The charge is highly variable and
reflects approximately 50 percent custom ore. The 24 installed Hirschoff
multi-hearth roasters are old and the source of considerable leakage.
Only 18 are normally in operation at a time. Offgases are treated in a
new, recently-installed electrostatic precipitator before being vented
to the atmosphere through a 544 ft stack. Vent gases from the roaster
load-out stations after cleanup in a baghouse are also directed into the
roaster offgas system. Calcines from the roasters are charged to 3
reverberatory furnaces. Reverberatory off gases pass through waste heat
boilers and a new electrostatic precipitator system before being vented
through the same 544 ft stack. Four of the five available Fierce-Smith
converters are usually in operation. Hoods of the Walter radiation
c°ollng type have been fitted recently to the converters but they are
n°t close fitting and air dilution is still appreciable. The off gases
are cleaned in electrostatic precipitators and then vented through a
separate 560 ft stack.
The location of the smelter in a valley limits the availability of
larid in the immediate vicinity. Utility services are fully loaded.
The smelter has no present S02 control system installation plans
ut Phelps Dodge management has stated that Arizona's ambient air standards
°r SO at Douglas would be met met through
1) a permanent cutback in smelter operation
2) the installation and operation of an intermittent
production curtailment system and
3) such operating adjustments as might thereafter be
shown to be necessary or desirable.
19
>:Ll-2 Weak S02 Streams
The offgases from the roasters, reverberatory furnaces and the
nv
Qnverters are all weak S02 gas streams as a result of the high degree
dilution in the various systems. A flow schematic is provided in
257
-------
Figure 12.11-1. With custom ores making up a significant part of the
smelter input, the smelter charge tends to be variable and this situation
might be expected to be reflected in variable operating conditions and
changing levels in gas flows and SO- contents as ore characteristics
change. The three gas streams have been characterized at the stack end
of the respective flue systems as shown in Table 12.11-1.
12.11.3 S02 Control Process Selection
The application of absorption-based SO,, control processes to the
entire gaseous output of this smelter—i.e., a total maximum gas flow of
approximately 700,000 SCFM with a total S02 loading of around 100,000
Ib/hr—is a rather overwhelming consideration. With all three offgas
systems having the same approximate low SO- content, there does not
appear to be any particular merit in attempting to use the S02 content
of any two of the streams to increase the S0? level of the third to
provide a satisfactory feed to a new, conventional sulfuric acid plant.
Although the roaster and reverberatory gas systems are independent and
presently share only the stack in common, it appears possible to combine
these two streams after their respective precipitators and install a
common gas conditioning and SO- absorption system. Similarly, although
there would be construction difficulties, the converter offgases could
be directed to a gas conditioning and S0~ absorption system before being
redirected back to the converter stack. These two absorption systems
could then supply a common S0? handling or regeneration section, sized
to handle the combined S02 rate and located in a convenient area of the
plant site, away from the congested smelting operational area itself.
In reviewing the available processes against the gas characteristics
and conditions at this smelter, the high oxygen levels of the gases,
together with the relatively dilute SO,, concentrations, suggest that the
Wellman-Lord, DMA/xylidine and ammonia processes would be penalized
substantially in terms of operating costs and/or higher capital investment.
The lime/limestone process for an S0? rate of approximately 100,000
Ib/hr would be associated with very large sludge rates and the limited
land availability around the Douglas smelter would seem to preclude this
process even without the potential operational problems which might be
expected with gas streams of high oxygen contents. The use of a direct
application sulfuric acid plant is clearly impractical. The magnesium
oxide process is associated with high energy demands and fairly large
258
-------
Ui
VO
PHELPS DODGE — DOUGLAS, ARIZONA
SMELTER FLOW SCHEMATIC
(181
ROASTERS
(24)
/ 260-290,000 SCFM \
1 1-8%SO2(AV) I
\ 380°F. /
N OPERATION AT TIME) / 145_156>0oo SCFM \ \
I 1-2% SO2 I \
\ AKff 1 \
REVERBS
(3)
CONVERTERS
v y \
(210-265.000 SCFM \
2-4% SO2 I
350°F. /
tSH • — ^
STACK 500 FT
410°F.
STACK
563 FT.
(4 IN OPERATION AT TIME)
FIGURE 12.11-1
-------
space requirements to accommodate the solids handling equipment, auxiliary
gas cleaning facilities and an auxiliary acid or elemental sulfur plant.
It does not appear to be an attractive match for this smelter. The two
control processes which might have some applicability to the Douglas
smelter are the double alkali and the citrate processes. However, the
former process still entails the disposal of impressive quantities of
throwaway sludge, and the task of trucking around the clock approximately
350,000 Ib/hr of sludge appears to undercut the practicability of this
approach. It is possible that a conveyor belt system to a depleted mine
site might provide an approach but the uncertainties involved suggest
that the only viable control is offered by the citrate process producing
elemental sulfur. The fact that only 1/2 Ib of sulfur is generated for
every pound of SCL removed and that this material can be readily stock-
piled is an impressive argument for the potential of this control process
in relation to the characteristics and special factors of the Douglas
smelter.
12.11.4 S02 Control Process Costs
The cost development for the citrate process will include two
separate, appropriately sized gas conditioning and S02 absorption systems
feeding a single SO^ regeneration and processing section. The maximum
gas flows and the average S02 rates provided in Table 12.11-1 will
provide the cost basis.
Special factors or conditions which modify the basic cost structures
are as follows:
1) The capital costs from Appendix A for gas cooling and
conditioning for the identified maximum gas flows for the
combined roaster/reverberatory gas streams and the
converter gas stream, have been modified by the factor
T«
0.6
R(t°F)
T R(600°F)
. •
where t°F = 410°- roaster/reverberatory gases
= 350°- converter gases.
2) The present congested flue configuration around the roaster/
reverberatory stack and the short flue length at the converter
stack will make the installation of the gas conditioning
and SO™ absorption systems with appropriate bypasses and
dampers at these locations extremely difficult. Factors
of 50 percent of the capital cost of the roaster/reverberatory
gas treating and absorption systems and 70 percent of the
capital cost of the converter gas treating and absorption
systems have been allowed to cover the additional costs involved.
260
-------
3) Additional surge capacity will be required at the central
SCL regeneration facility to handle the pregnant citrate
liquor returns from the two absorption systems. An
allowance of $100,000 total capital has been provided
and added to the S02 regeneration facility capital cost.
4) Additional electrical substations and water treatment
facilities to support the makeup water requirements of
the scrubbing systems and the steam generation units
associated with the citrate process have been provided,
with the estimated costs being based on the unit
relationships established for the gas conditioning, S0?
absorption and S02 handling sections of the citrate process
and the cost curves in Appendix F.
5) To cover the additional costs associated with implementing
a major process change in an operating plant, an estimate
of 15 percent of the total boundary limit cost for the
control process has been allowed.
The capital and operating cost structures for the citrate process
provided in Tables 12.11-2 and 12.11-3, respectively. The individual
computation sheets are provided under the Phelps-Dodge-Douglas smelter
in Appendix G.
261
-------
Table 12.11-1. PHELPS-DODGE CORPORATION - DOUGLAS, ARIZONA
CHARACTERIZATION OF WEAK S02 GAS STREAMS
Stream
Roasters (24)
(18 in operation
at a time)
Reverbs (3)
Converters (5)
(4 in operation
at a time)
Volume
SCFM
260-290,000
145-156,000
210-265,000
Av. S02 Content
Sizing basis (Ibs/hr)
51,100
19,000
43,000
Equivalent Av.
S02 Content (%)
1.8
1.25
1.8
% Oxygen(2)
14-18
10-15
12-18
Temp.°F
380
450
350
to
• to
Notes.
(1) SO., loading in all cases approximately 0.1%.
Particulates less than 0.01 grains/SCF for the
roaster and reverberatory streams and less than
0.02 grains/SCF for the converter stream.
(2) Oxygen content of gas streams estimated.
-------
Table 12.11-2. PHELPS DODGE - DOUGLAS, ARIZONA
CAPITAL COSTS FOR SO2 CONTROL PROCESSES ON WEAK S02 SMELTER OFF-GASES
ITEM
1. Boundary limits retrofitted
primary system
2. Auxiliary Plant:
(a) H2SO, plant with dry gas
cleaning
TOTAL Boundary Limit Cost
3. Total Service Costs
4. Associated Investment Costs
(a) new ESP
(b) General site-related costs
Total Capital Investment
Capital cost/annual tons SO^
Removed
Primary S02 Control Method
Citrate Process
(95% Removal)
39,150,000
N/A
39,150,000
1,900,000
5,870,000
$46,920,000
$107
COMMENTS
to
-------
Table 12.11-3. PHELPS DODGE - DOUGLAS, ARIZONA
ANNUAL OPERATING COSTS AND TOTAL NET ANNUALIZED COSTS
ITEM
Primary SO- Control Method
Citrate Process
(95% Removal)
T
COMMENTS
CT>
1. Gas conditioning and S02
absorption
2. S02 handling
3. Labor
4. Maintenance, Ins. & Taxes
5. Total Annual Operating Cost
6. Auxiliary Plant Total
Operating Costs
(a) Sulfuric acid plant
7. Total Annual Operating Costs
8. Annualized Capital Cost
9. Total Net Annualized Cost
10. Cost/year/ton SO- removed
11. Cost/lb copper
1,062,000
8,914,000
561,000
2,873,000
13,410,000
N/A
$13,410,000
$ 6,170,000
$11,642,000
$27
4.48c/lb
Based on smelter's present operating level of 2000 TPD concentrate, 20% copper 130,000 TPY.
net-back to smelter for sulfur produced.
Zero
-------
12.12 PHELPS DODGE CORPORATION -- MORENCI, ARIZONA
12.12.1 Smelter Characteristics
The Morenci smelter began operations in 1943. Concentrate capacity
is over 2000 TPD. Approximately 35 percent of the total plant feed is
processed through a fluid bed roaster with the balance feeding directly
up to 3 of the 5 available reverberatory furnaces. Four of the re-
verberatories are old units and are approximately the same size, but the
fifth unit, commissioned in 1974, is a much larger furnace. A total of
nine Fierce-Smith converters are available, and under normal scheduling
with seven operating, they provide a relatively even offgas flow although
S0,j concentration does vary. All converters are fitted with close-
fitting hoods with appropriate dampers to provide shutoff when not
blowing and volume and system balancing control.
Offgases from the fluid bed roaster are mechanically cleaned in a
series of cyclones before passing to the gas cleaning system of an old
single contact sulfuric acid plant originally rated at 750 TPD but
actually with a present maximum capacity of around 600 TPD acid. The
roaster gas S0? concentration of approximately 12-15 percent is diluted
to about 6.5 percent S09 before being introduced into the acid plant.
Reverberatory offgases pass through waste heat boilers and electro-
static precipitators before being vented out of a 605 ft stack.
Gases from the converters pass through coolers and electrostatic
Precipitators to the gas cleaning plant and then through an extended
flue system to a dual train, single contact sulfuric acid plant rated at
2500 TPD of acid (237,000 SCFM at 5.5 percent S02>. Most of the acid
Presently produced at Morenci must be disposed of by neutralization.
^o tailings leach/lime neutralization modules have been constructed
With a total capacity of about 1800 TPD of acid.
The smelter is extremely congested and available space around the
teverberatory operating area is minimal. Five 150 ft diameter thickeners
are located in close proximity to the smelter operating area. Water
8uPPly is limited.
2>12.2 Weak S02 Streams
The only weak S02 streams in this smelter are the offgases from the
*everberatories. The furnaces are operated at a slight negative pressure
the usual operations of routine charging contribute to air infiltration.
265
-------
Routine "soot" blowing of the waste heat boilers and leakage at the
electrostatic precipitators also contribute to the considerable degree
of air dilution experienced in this system. Depending on the line-up of
the available reverberatory furnaces, gas flows in the reverberatory
flue gas systems at the stack have been estimated to range from 340-
500,000 SCFM. A flow schematic is provided in Figure 12.12-1. the
reverberatory gas stream has been characterized at the stack as shown in
Table 12.12-1.
12.12.3 SC)2 Control Process Selection
Although the high oxygen ratio in the gas stream is unfavorable in
both operational and economic terms to most of the candidate SO^ ab-
sorption based control processes, other considerations are equally
important at this smelter. Morenci is already a major producer of
sulfuric acid, and while the general market for acid and the corresponding
price has moved upward markedly over the past year, it is apparent that
present geographical conditions limit the available market and Morenci
has found it necessary to neutralize a significant proportion of their
existing acid production. An S0~ control system on the reverberatory
gases with the production of another potential 700-800 TPD of acid does
not appear to be a very attractive option either now or in the longer
term, even if the demand for acid in Arizona increases appreciably.
The majority of the regenerable processes do provide the S02 end-
use options of both elemental sulfur and liquid SO™ in addition to
sulfuric acid, but the potential economic penalties associated with the
high oxygen levels in the gas stream and related high oxidation rates in
the Wellman-Lord, xylidine and ammonia processes appear to preclude
serious consideration of these three systems. The magnesium oxide
process produces only a 15 percent stream of regenerated S0~ and the
elemental sulfur route is unattractive when evaluated in terms of natural
gas (reductant) requirements and related costs. The citrate process
with its limited oxidation potential and direct production of elemental
sulfur appears to be the most attractive process although the system is
still in the developmental phase. Of the two throwaway processes only
the double alkali appears to offer a reasonable approach under the
Morenci conditions. The direct application of a sulfuric acid plant to
the reverberatory gases is obviously unrealistic in view of the comments
made above and the very high energy demand of this approach.
266
-------
PHELPS DODGE — MORENCI, ARIZONA
SMELTER FLOW SCHEMATIC
S3
I
I
I
I
i
I
i
R|DASTERJ
(FLUID BED)
I I
I I
REVERBS
(5)
-u-
CYCLONES
I 1
J GAS I
I CLEANING
1
I
u _
1
r
i
i
i
i
i
i_
i
ACID j
PLANT (S.C.) 1
600 TPD MAX. \
I
STACK
(38-42,000 SCFM
6.5% SO2
340-500,000 SCFM
1.0% SO2 (AV.)
450°F.
WHB
ESP
STACK
605 FT.
180-237,000 SCFM A
3.5%-7.0% SO2 J
I T
I I
I CONVERTERS I
! O) I-
j
i
I
ESP
i !
i 1
i 1
--J GAS H
j CLEANING I
r — — i
} ACID PLANT !
1 (S.C.) L
| DUAL TRAIN. {
STACK
(7 IN OPERATION AT TIME)
2500 TPD
[237,000 SCFM @ 5.5% SO21
FIGURE 12.12-1
-------
Thus the only SCL control processes judged to be applicable to the
Morenci reverberatory gas stream are the double alkali and citrate
processes.
12.12.4 S02 Control Process Costs
The special factors or conditions which modify the basic cost
structures for this smelter are as follows:
1) The temperature of 450°F at the stack end will reduce the
capital costs from Appendix A for the necessary gas cooling
and conditioning. The modification factor is
T
o
R(450°F)
T
o
R(600°F)
0.6
2) Retrofitting a gas conditioning and absorption system to the
reverberatory gas duct system and accommodating the equipment
within the extremely congested area in the immediate vicinity
of the duct system appear to be major problems. There are
6 acid storage tanks located parallel to and almost under
the flue system with 2 adjacent railroad spins. Special
approaches towards equipment configuration and connecting
ducting would seem mandatory. A factor of 70 percent of the
capital cost of the gas conditioning and S02 absorption
systems has been allowed to cover the additional costs
3) To provide uniform processing rates in the S02 regeneration
section, additional surge capacity to handle the S02 rate
variations in the gas system will be required at the S02
absorption section. An allowance of 5 percent of the S02
absorption section cost has been provided.
4) Additional electrical substations and water treatment facilic
to support the makeup water requirements have been provided»
with the estimated costs being based on the unit relationship
established for the appropriate systems and the cost curves
in Appendix F.
5) Any major process changes in a plant as congested as Morenci
will incur substantial additional costs related to coordinat
and scheduling difficulties, special safety requirements, et
An allowance of 20 percent of the total boundary limit cost
for each of the selected processes has been provided.
The capital and operating cost structures for the double alkali a
"
citrate processes are provided in Tables 12.12-2 and 12.12-3, respec
Individual computation sheets are provided under the Phelps Dodge-
Morenci smelter section in Appendix G.
268
-------
Table 12.12-1. PHELPS DODGE CORPORATION - MORENCI, ARIZONA
CHARACTERIZATION OF WEAK S02 REVERBERATORY GAS STREAMS.
Stream
Reverb. (1)
Volume
SCFM
340-500,000
Av. S02 Content
Sizing basis (Ibs/hr)
43,000
Equivalent Av.
S02 Content (%)
1.0
% Oxygen(2)
12-15
o
Temp . F
450
VO
Notes. (1) SO- reported as 0.1% at furnace uptake.
(2) Oxygen content of gas stream estimated.
-------
Table 12i2-2. PHELPS DODGE CORPORATION - MORENCI, ARIZONA
CAPITAL COSTS FOR S02 CONTROL PROCESSES ON REVERBERATORY FURNACE OFF-GASES
ITEM
1. Boundary limits retrofitted
primary system
2. Auxiliary Plant
(a) H2S04 plant with dry
gas cleaning
TOTAL Boundary Limit Cost
3. Total Service Costs
4. Associated Investment Costs
(a) new ESP
(b) General site-related
costs
Total Capital Investment
Capital cost/annual ton S02
Removed
Primary S02 Control Method
Double Alkali
(95% Removal)
21,880,000
N/A
21,880,000
1,120,000
4,370,000
27,370,000
$164
Citrate Process
(95% Removal)
25,150,000
N/A
25,150,000
1,360,000
5,000,000
31,510,000
$189
COMMENTS
to
•-4
O
-------
Table 12.12-3. PHELPS DODGE CORPORATION - MORENCI, ARIZONA
ANNUAL OPERATING COSTS AND TOTAL NET ANNUALIZED COSTS
ITEM
1. Gas conditioning and SO-
absorption
2. SO' handling
3 . Labor
4. Maintenance, ins. & taxes
5. Total Annual Operating Cost
6. Auxiliary Plant Total
Operating Costs
(a) Sulfuric acid plant
7. Total Annual Operating Cost
8. Annualized Capital Cost
9. Total Net Annualized Cost
10. Cost/year/ton SO- Removed
11. Cost/lb copper &'
Primary S02 Control Method
Double Alkali
(95% Removal)
713,000
7,161,000
210,000
1,622,000
9,706,000
N/A
9,706,000
3,599,000
7,770,000
$47
2.22c/lb
Citrate Process
(95% Removal)
800,000
3,387,000
333,000
1,893,000
6,413,000
N/A
6,413,000
4,144,000
6,470,000
$39
1.85C/lb
COMMENT
(1)
Based on smelter rated capacity of 2200 TPD concentrate , 24% copper operating 340 days/year.
Annual production 175,000 TPY. Zero net-back to smelter for sulfur produced.
-------
272
-------
12.13 WHITE PINE COPPER COMPANY — WHITE PINE, MICHIGAN
12.13.1 Smelter Characteristics
The White Pine Copper Company smelter is one of the newer U.S.
copper smelters, having begun operations in the mid-fifties. Siting of
the plant was chosen taking atmospheric dispersion characteristics into
consideration. Green concentrate mixed with coal is charged to the two
reverberatory furnaces by conveyor belt with charging on each furnace
occupying about 9 hours out of every 24. Matte from the reverberatories
is converted in two Fierce-Smith converters with only one on line at a
time. Gases from the reverberatory furnaces pass through waste heat
boilers to a common flue and are cleaned in an electrostatic precipi-
tator before being vented up a 504 ft stack. Considerable dilution of
these gases takes place through the flue although the system has been
recently renovated. The waste heat boilers are "soot" blown each shift
for about one hour using steam. Additional steam is also introduced
continuously into the electrostatic precipitator. The converter gases
are also vented directly to the 504 ft stack but without special parti-
culate removal treatment.
It must be noted that the smelting operation at the smelter is not
typical of domestic primary copper smelters. The copper values of the
°*e-body mineralogy are present exclusively as chalcocite and native
copper , and reverberatory processing of the green concentrate as is
*0uld produce a difficult to handle matte containing over 75 percent
c°Pper. The copper content is reduced to a manageable level (about 65
percent) by including purchased pyrite with the reverberatory furnace
charge. The 65 percent matte requires only a brief slag blow (30 minutes
°tal for both the first and second slag blow) to go to white metal.
Th
ne total converter operating cycle is 4 hours with a total elapsed
time of 10 1/2 hours. With only 2 converters available, gas flows
t0ln this operation are thus intermittent and limited to about 8-9 hours
ln any one day.
This smelter under its current operating mode has demonstrated its
0lnPliance with ambient air standards through 1980. It has been indi-
ated that approaches towards emission level reductions after this date
Will ,
A tocus on metallurgical and related processing considerations.
273
-------
The smelter is located in north Michigan on Lake Superior with
adequate land availability. The surrounding area is heavily treed.
12.13.2 . Weak SO,, Streams
The offgases from both reverberatory furnaces contain a very low
level of S0» even with the inclusion of pyrite in the reverbera.tory
charge. Air infiltration in the reverberatory flue system further
reduces the SCL content to marginal values. Figure 12.13-1 provides a
flow schematic of the gas system and indicates the gas flows and S0~
concentrations under the usual operating mode.
With only two converters, gas flows in the converter flue system
are intermittent, with flows ranging between approximately 99,000 SCFM
for 20 minutes duration during the first slag blow to 142,000 SCFM for
10 minutes during the second slag blow. The copper blow duration is
about 3 1/2 hours with an associated gas flow of approximately 106,000
SCFM. Sulfur dioxide contents vary from 1 percent during the first slag
blow to 2.5 percent during the second slag blow to 4.5 percent during
the copper blow. Total SO,, generation during a 24 hour period has been
computed on the basis of the 10 1/2 hour elapsed time converter cycle
and the 4 hour operational cycle.
Although the converter gas streams in this smelter do not fall
within the accepted definition of "weak" S02 gas streams, they present a
special problem because of their intermittent nature. They have been
characterized in terms of gas flow and average S02 content together with
the reverberatory gas stream in Table 12.13-1.
12.13.3 S00 Control Process Selection
z
In view of this smelter's present compliance with ambient air
standards, its location, and the marginal S02 level of the reverberatory
gases, specific application of an SO,, control process to the rever-
beratory gas stream does not appear to be warranted. The converter
system, however, can be controlled with an absorption-based process if
sufficient surge capacity is provided to operate the regeneration section
continuously and the gas conditioning and SO- absorption system is sized
to the maximum expected gas flow. It is noted that the maximum gas flow
of 142,000 SCFM which is associated with the second slag blow is only of
10 minutes duration and that a sizing level of 110,000 SCFM would accommo-
date the complete gas flow rate for at least 96 percent of the time.
274
-------
— WHITE PINE, MICHIGAN
SMELTER FLOW SCHEMATIC
160-185,000 SCFNT
0.17% SO2
12% O2
500°F
REVERB NO. 2
WHB
to
-vl
REVERB NO. 1
WHB
CONVERTER
(2)
INTERMITTENT FLOW (9-10 HRS APPROX/DAY)
STACK
504 FT
99-142,000 SCFM
4.5% SO2
17.0% O2
\550°F
(1 ON USE AT
TIME)
NOTE: (1)
FEED EXCLUSIVELY CHALCOCITE AND NATIVE COPPER
PYRITE ADDED TO CONTROL MATTE COPPER CONTENT
(2) BASED ON ACTUAL TEST DATA TAKEN IN MAY 1975
FIGURE 12.13-1
-------
The very high oxygen level in the converter gas stream will limit
applicable SCL control processes to those least affected by high oxida-
tion levels. Of the regenerable processes, only the magnesium oxide and
citrate processes appear to be viable candidates, although as noted
previously, there is some uncertainty that sulfate formation in the MgO
process is self-limiting at around 15 percent. Without this provision,
purging and secondary treatment of the MgS04 with lime [Ca(OH>2] will be
necessary to recover the magnesium. The double alkali throwaway process
also appears to be applicable. Although high oxidation levels of the
sodium sulfite/bisulfite will occur, the lime regeneration step plus
appropriate softening should provide an acceptable process.
12.13.4 S00 Control Process Costs
/ "
The special conditions associated with the White Pine Smelter which
modify the cost structures of the selected processes are as follows:
1) The average temperature of the converter gas stream at the
stack is 550°F and the capital costs from Appendix A for gas
cooling and conditioning have been modified by the factor
T R(550°F)
T R(600°F)
0.6
2) Retrofitting the gas conditioning and S02 absorption sections
to the existing flues and providing bypasses and appropriate
dampers incur additional costs, and these have been allowed
for on the basis of 15 percent of the capital investment of
these two sections.
3) To provide continuous operation of the S02 regeneration
sections of the control processes, large surge capacity after
the absorption system will be necessary. Actual absorption
will occupy only 9-10 hours of each day. An additional 7
percent of the capital investment of the SO^ absorption
section has been provided.
4) Additional service facilities required will include an
electrical substation, but specific water treatment facilities
should not be necessary. Capital cost curves are provided in
in Appendix F.
5) An auxiliary sulfuric acid plant taking an 8 percent S02
gas stream has been provided with the magnesium oxide process.
Capital and operating costs of this plant based on the
S09 hourly rate of the magnesium oxide process itself are
provided in Appendix C.
276
-------
6) An estimate of 20 percent of the total boundary limit
costs for each process has been allowed to cover the
additional costs usually associated with the
implementation of major process changes in an ongoing
operating plant.
The capital and operating cost structures for the three processes
identified for the White Pine Smelter—double alkali, magnesium oxide,
and citrate processes—are provided in Tables 12.13-2 and 12.13-3,
respectively. Individual computation sheets for each process are provided
under the appropriate smelter name in Appendix G.
277
-------
TABLE 12.13-1. WHITE PINE COPPER COMPANY - WHITE PINE, MICHIGAN
CHARACTERIZATION OF S02 GAS STREAMS
to
~J
CO
Stream
Reverb. (2)
Converter
Volume
SCFM
160,000-
185,000
99,000-
142,000
(Predominant
flow rate
106,400)
Av. SO 2 Content
Sizing basis (Ibs/hr)
2,800
176,300 lb f or 4hr cyd
On basis of 2.5 cycles/
day, av. S02 rate for
24 hours =18, 500 Ib/hr
Equivalent Av.
SO- Cont. (%)
0.17
e (A. 5)
(2)
% Oxygen
12
17
Temp. °F
500
550
Notes. (1) No data on SO- or particulate loading.
(2) Oxygen levels determined during test programs.
-------
Table 12.13-2. WHITE PINE COPPER COMPANY - WHITE PINE, MICHIGAN
CAPITAL COSTS FOR SO CONTROL PROCESSES ON CONVERTER OFF-GASES
ITEM
1. Boundary limits retrofitted
primary system
2. Auxiliary Plant:
(a) H2S04 plant with dry
gas cleaning
TOTAL Boundary Limit Cost
3. Total Service Costs
4. Associated Investment Costs
(a) new ESP
(b) General site-related
costs
Total Capital Investment
Capital cost/annual ton. S0_
Removed
Primary SO^ Control Method
Double Alkali
(95% Removal)
9,420,000
N/A
9,420,000
320,000
1,880,000
11,620,000
$163
Magnesium Oxide
(90% Removal)
9,120,000
3,750,000
12,870,000
350,000
2,570,000
15,790,000
$240
Citrate
(95% Removal)
10,500,000
N/A
10,500,000
330,000
2,100,000
12,930,000
$181
COMMENTS
N>
-------
Table 12.13-3. WHITE PINE COPPER COMPANY - WHITE PINE, MICHIGAN
ANNUAL OPERATING COSTS AND TOTAL NET ANNUALIZED COSTS
ITEM
1. Gas conditioning and SO,,
absorption
2. SO £ handling
3. Labor
4. Maintenance, Ins. & Taxes
5. Total Annual Operating Cost
6. Auxiliary Plant Total
Operating Costs >,.
(a) Sulfuric acid plant
7. Total Annual Operating Cost
8. Annualized Capital Cost
9. Total Net Annualized Cost
10. Cost/year/ton SO. removed
11. Cost/lb copper (2)
Primary S0_ Control Method
Double Alkali
(95% Removal)
176,000
3,137,000
210,000
686,000
4,209,000
N/A
4,209,000
1,528,000
3,345,000
$47
Magnesium Oxide
(90% Removal)
200,000
2,039,000
315,000
757,000
3,311,000
450,000
3,761,000
2,076,000
3,527,000
$54
2.35c/lb
Citrate
(95% Removal)
200,000
1,453,000
333,000
770,000
2,756,000
N/A
2,756,000
1,700,000
2,720,000
$38
1.81C/lb
COMMENTS
to
00
o
Overall SO,, removal efficiency with H SO, plant is 87.3%.
(1)
(2)
Based' on annual production of 75,000 TPY. Zero net~back to smelter for acid or
sulfur produced.
-------
APPENDIX A
GAS CLEANING AND CONDITIONING
In establishing costs for a gas cleaning and conditioning system
prior to an absorption-based S02 control process, an attempt has been
made to recognize the actual demands of the processes themselves.
The pretreatment requirements of feed gases to sulfuric acid plants
have been well defined in terms of both residual concentrations of
contaminants to minimize catalyst plugging and deterioration, and the
process equipment to accomplish this standard. Acid plant installations
in smelter environments are characterized by extensive gas treatment
facilities, which constitute a significant proportion of the total
investment.
The use of weak SO,, streams as input to sulfuric acid plants directs
even greater emphasis to the gas pretreatment section since additional
cooling or even refrigeration is necessary to reduce the moisture content
°f the gas feed as part of maintaining the required water balance.
Because of the interrelationship between conditioning requirements and
SO- content of weak streams feeding sulfuric acid plants, gas conditioning
for sulfuric acid plants has been treated as part of the recovery process
itself and is not covered specifically in this section.
As discussed in Section 2.0 both regenerable and throwaway S02
control processes require a cooled and humidified input gas to reduce
excessive evaporation of the aqueous S02 absorbent with attendant scaling
a^d related problems. Temperatures between 130° and 150°F for most
Processes appear acceptable, although certain processes, namely the NH.,
and citrate systems, appear to require lower temperatures of around 90°
and 125"F, respectively. Dry gas cleaning on the smelter gas streams
accompllshed by cyclones, balloon flues and electrostatic precipitators
"as the capability of reducing the particulate and fume material to
residual levels of 0.1 to 0.2 gr/SCF. The presence of this residual
material may introduce certain problems to the S02 absorption and recovery
clrcuits such as:
1) buildup of solids in closed loop systems
2) carryover into the final product
3) increased oxidation of active salts by catalysis
or direct interaction.
281
-------
4) equipment corrosion if halides and mercury compounds
are present.
The wet cleaning system should be designed to minimize the effect of
these contributions on the performance and economics of each specific
control process.
A flow sheet for a typical gas cleaning system for a regenerative-
SCL control system which provides the required steps of cooling and
humidification and removal of particulate and fume material is shown in
Figure A-l. Inlet conditions are taken as 600°F temperature, 0.03
percent SC>3, a particulate loading of 0.1 to 0.2 gr/SCF and are assumed
to be representative of a copper smelter reverberatory furnace gas
stream after the usual dry gas cleaning. Cooling and humidification is
commonly accomplished in a spray tower or impingement column with re-
circulation of the scrubbing water although the use of high energy
venturi absorbers is highly effective where energy is readily available.
The absorption of SO., by the recirculated scrubbing water produces a
weak H^SO, acid solution which can theoretically reach a concentration
of 50 to 60 percent sulfuric acid, but where halides are present in the
gas stream, acid strength is limited to 5 to 10 percent by discarding a
portion of the recirculating stream usually to a neutralization section
and using fresh water makeup. Additional cooling from the 150°F saturation
temperature in the scrubbing tower to 130°F or lower, if necessary, is
provided in a second packed tower using cooled, recirculated weak acid.
This recirculated stream is cooled in graphite or stainless steel heat
exchangers with water recirculated through a water cooling tower loop.
The cooled gases leaving this column still contain both some fine
particulates and sulfuric acid mist aerosols from the S0~ entering with
the feed gas. Treatment in wet ESP's and/or high efficiency demisters
is the final step in the conditioning sequence.
For the non-regenerable or "throwaway" S0~ control processes, the
gas pretreatment requirements are satisfied by scrubbing and cooling the
feed gas only, and the system requirements accordingly are less demanding.
2 3
Capital costs have been developed from reported data ' for systems
similar to that depicted by Figure A-l with appropriate adjustment for
equipment differences, escalation, indirects, etc.; and a cost versus
gas flow rate curve is provided in Figure A-2. Operating costs have
282
-------
150°F
SPRAY
OR _
IMPINGEMENT
TOWER
r-o
oo
SMELTER
GAS -
600° F
LIMESTONE
730°F
COOLING
COLUMr
NEUTRALIZATION
TO POND
TO
S0
ADSORPTION
ELECTROSTATIC
MIST
PRECIPITATOR
COOLING TOWER
GAS COOLING & CONDITIONING FOR REGENERATIVE S0? ABSORPTION SYSTEMS
FIGURE A-1
-------
been calculated using reported or developed values for power and water
requirements with limestone quantities based on a 0.03 percent input
level of SO.,.
Usage factors and unit cost data together with maintenance and
labor factors are presented in Table A-l.
Capital and total annual costs have been calculated for four gas
flow rates varying from 70,000 to 300,000 SCFM for each of the two gas
conditioning systems and these are presented in Tables A-2 and A-3,
respectively.
REFERENCES
1. Donovan, J.R. and P.J. Stuber, "Sulfuric Acid Production from Ore
Roaster Gases," J. Metals 19 (11) 45-50 (November, 1967).
2. Industrial Gas Cleaning Institute, Inc., "Air Pollution Control
Technology and Costs in Nine Selected Areas," PB-222-746, September
1972 pp. 294-298.
3. Bureau of Mines RI 7957, "Cost of Producing Copper from Chalcopyrite
Concentrate as Related to S0? Emission Abatement," p. 44.
284
-------
Table A-l. GAS CLEANING AND CONDITIONING UNIT USAGE AND COST DATA
A. Chemicals & Utilities
Limestone: Regenerable System
(Throwaway System)
Power: Regenerable System
(Throwaway System)
Make-up Water
B. Operating Labor &
Maintenance
Labor3
Maintenance
Taxes and Insurance
Basis
0.09 Ib/M SCFM1
0.07 Ib/M SCFM
7,0 KW/M SCFM
5.25 KW/M SCFM
3.0 gal/M SCFM2
*5 man/shift
< 75,000 SCFM
3/4 man/shift
> 75, 000 SCFM
5% TCI/year
2 1/2% TCI/year
Unit Cost
$8.0/ton
$0.015/KWHr
$0.30/M gal
$8.0/hr
c- Fixed Charges 13.15% TCI/year
Based on Capital Recovery Factor using 10% interest over 15 year life
fiased on 0.03% S0~ in in-coming gas.
2
Delude losses from cooling tower.
3
Labor costs taken over 365 days.
285
-------
GAS CONDITIONING fie CLEANING
TOTAL CAPITAL INVESTMENT COSTS
& TOTAL ANNUAL DIRECT OPERATING COSTS
CAPITAL INVESTMENT
(REGENERABLE
SYSTEMS)
$1.0 Million
CAPITAL INVESTMENT
(THROWAWAY SYSTEMS)
ANNUAL OPERATING
COSTS
(REGENERABLE
SYSTEMS)
ANNUAL OPERATING
COSTS (THROWAWAY
SYSTEMS)
100,000 SCFM
GAS FLOW RATE-SCFM
286
-------
Table A-2. GAS CLEANING AND CONDITIONING FOR REGENERABLE SO2 CONTROL PROCESSES
CAPITAL AND TOTAL ANNUAL COSTS
N>
00
TOTAL CAPITAL INVESTMENT
Annual Cost
A. Direct Operating
1. Power
2. Water
3. Limestone
4. Labor
5. Maintenance
6. Taxes & Insurance
TOTAL ANNUAL DIRECT
OPERATING COST
B. Annual Capitalized
Cost
C. Net Total
Annualized Cost*
$
$/SCFM
$/SCFM
$/SCFM
Gas Flow Rate - SCFM
70,000
1,800,000
26
60,000
30,800
12,300
35,000
90,000
45,000
273,100
3,9
236,700
321,100
4.6
100,000
2,275,000
23
85,700
44,000
17,600
52,600
113,800
56,900
370,600
3.7
299,200
419,100
4.2
*Based on Corporate Tax Rate of 48%. •-
200,000
3,550,000
18
171,400
88,000
35,200
52,600
177,500
88,800
613,500
3,1
466,800
672,200
3.4
300,000
4,650,000
16
257,100
132,000
52,800
52,600
232,500
116,300
843,300
2.8
611,500
740,000
2,5
-------
Table A-3. GAS CLEANING AND CONDITIONING FOR "THROWAWAT'SO- CONTROL PROCESSES
CAPITAL AND TOTAL ANNUAL COSTS
TOTAL CAPITAL INVESTMENT
Annual Cost
A. Direct Operating
1. Power
2. Water
3. Limestone
4 . Labor
5. Maintenance
6. Taxes & Insurance
TOTAL ANNUAL DIRECT
OPERATING COST
B. Annual Capitalized
Cost
C. Net Total
Annualized Cost*
$
$/SCFM
$/SCFM
$/SCFM
70,000
1,210,000
17
44,900
30,800
9,600
35,000
60,500
30,300
211,100
3.0
159,100
230,200
3.3
Gas Flow Rate
100,000
1,500,000
15
64,200
44,000
13,700
52,600
75,000
37,500
287,000
2.9
197,300
298,500
3.0
- SCFM
200,000
2,250,000
11
128,400
88,000
27,400
52,600
112,500
56,300
465,200
2.3
295,900
465,800
2.3
300,000
2,900,000
10
192,300
132,200
41,100
52,600
145,000
72,500
635,700
2.1
381,400
619,200
2.1
to
00
00
*Based on Corporate Tax Rate of 48%.
-------
APPENDIX B
S02 ABSORPTION COSTS
The absorption section of absorption-based SCL control systems
generally includes the gas scrubbing unit itself, the absorbent re-
circulation loop including hold-up and storage, and a suitable mist
separator to prevent carryover of the absorbent to the stack. Certain
processes, however, may require additional equipment to satisfy this
special requirement. The fan or blower necessary to move the gas through
both the conditioning and absorption sections may be located before the
conditioning section or after the SC^ absorber itself. Gas characteristics
flow rate, temperature, overall economics, may all exercise an influence
on a specific application. For this study, it is assumed that the blower
°r fan is located before the gas conditioning section, but as noted in
Section 2, the capital charges are allocated entirely to the S02 absorp-
tion section with the power operating costs being prorated over both
Actions based on expected pressure drop in each section.
Of the candidate S02 control processes under review, three are
characterized by a very specific S02 absorption step—sulfuric acid,
Dimethylaniline and ammonia scrubbing—which directly influences the
associated cost structure. The other processes, however, utilize either
8lurry or clear solution scrubbing; and while there is a variety in the
absorption equipment being used by these processes, the effect on over-
all capital costs does not appear to be significant. Slurry scrubbing,
because of the higher L/G ratios used and the effect on auxiliary equip-
"^nt, tends to incur capital costs somewhat higher than those for clear
a°lution scrubbing.
Cost studies for S00 control processes applied to power plants have
v 12 3
Deen reported in some detail by TVA, ' and Catalytic Incorporated.
°ther order-of-magnitude estimates for various S02 control processes
ave also been reported in trade journals and papers presented at various
8ymPosia. These data have been reviewed and used with appropriate
°dification, adjustment and escalation to develop the general capital
°.8t curves provided in Figure B-l. Escalation and indirect loading
a°tors used are the same as those specified in Section 2.
289
-------
A capital cost curve for the ammonia process has been developed
separately, but is represented in Figure 10-5. The dimethylanine process
absorption system is such an integral part of the control process itself
that the costs have not been separately identified.
Since operating costs are more clearly identified with the specific
control processes, a generalized approach has not been taken and the
costs have been developed as part of the effort for each particular
control process.
REFERENCES
1. G.G. McGlamery, et al. Tennessee Valley Authority, Conceptual
Design and Cost study, Sulfur Oxide Removal for Power Plant Stack
Gas (Magnesia Scrubbing-Regeneration), PB-222-509, May 1973.
2. G.G. McGlamery and R.C. Torstrick, Tennessee Valley Authority,
"Cost Comparisons of Flue Gas Desulfurization System", Paper presented
at Flue Gas Desulfurization Symposium, Atlanta, Georgia, November
1974.
3. E.L. Calvin, Catalytic, Inc., "A Process Estimate for Limestone
Slurry Scrubbing of Flue Gas", EPA-R2-73-1489, January, 1973.
290
-------
SO2 ABSORPTION SYSTEMS
TOTAL CAPITAL INVESTMENT COSTS
2-STAGE LIMESTONE
SCRUBBING
SLURRY SCRUBBING
SOLUTION SCRUBBING
Costs: Mid 1974
100,000 SCFM
GAS FLOW RATE-SCFM
291
FIGURE B-1
-------
APPENDIX C
SULFURIC ACID PLANT
Auxiliary sulfuric acid plants coupled to regenerable SO,, control
processes usually include only the acid-making section together with
appropriate gas drying and acid cooling and storage facilities. An
exception to this situation is posed by the magnesium oxide process
which regenerates the active absorbent and releases the captured S02 via
a high temperature calcination process. The resulting gas stream con-
taining about 15 percent S02 is both hot and "dirty" with entrained
particulates. Although the conventional approach is to cool and clean
these gases via water scrubbing, in view of today's concern with energy
utilization and conservation, it is both feasible and practical to
consider using a waste heat boiler for heat recovery and cooling, air
dilution to provide an approximately 8 percent S02 stream at 400-450°F,
and high temperature bag filters to accomplish particulate removal
before introduction to the acid plant drying tower,
Capital and operating costs have been developed for both situations,
i.e.,
1) auxiliary acid plant only with no gas cooling and cleaning
facilities
2) auxiliary acid plant with dry gas cleaning facilities (heat
recovery facilities have been included with the primary
S07 control system) .
Capital and operating costs for sulfuric acid plants have appeared
in a number of published reports and technical articles over the past 5
to 6 years. As with most published cost data, the bases and conditions
upon which the estimates have been made are not well defined, and it is
difficult to compare or evaluate such data. A number of formal references,
together with peripheral cost data gleaned from a range of articles
which included references to acid plant costs, were used to develop
costs for an acid section only, handling an 8 percent SO^ feed gas, and
for an acid section with conventional wet scrubbing with appropriate
allowances for engineering and contractor fees and storage costs. The
costs of dry gas cleaning facilities were developed from information
provided in the TVA conceptual design and cost study on the magnesium
oxide scrubbing process applied to power plant stack gases.
292
-------
Operating costs for an auxiliary acid plant only and for an auxiliary
acid plant provided with dry gas cleaning have been determined on the
basis of the usage and cost data provided in Table C-l. Figure C-l
provides both capital and direct annual operating costs for these two
operations over a range of SO,, values. These SO., values are based on an
acid plant efficiency of 97 percent and have been adjusted to represent
the S02 rate from the regeneration section of the primary control system
itself. To provide ready matching of capital and operating costs with
the primary control systems in terms of various gas flows at 1 percent
S0?, Tables C-2 and C-3 have been prepared on the basis of the gas flow
and S09 content to the primary SO,, control process.
REFERENCES
1. "The Production and Marketing of Sulfuric Acid from the Magnesium
Oxide Flue Gas Desulfurization Process," by Zonis, Hoist, Olmsted
and Cunningham, Essex Chemical Corp. Presented at EPA Flue Gas
Desulfurization Symposium, Atlanta, Georgia, November 1974.
2. "Economics of Sulfuric Acid Manufacture," J.M. Connor, Chemical
Engr. Progress, November 1968 (59-65) Vol. 64, No.11.
3. "The Impact of Air Pollution Abatement on the Copper Industry. An
Engineering Economic Analysis Related to Sulfur Oxide Recovery," by
Fluor Utah Engineers and Constructors, Inc. for Kennecott Copper
Corp, PB-208-293, April 1971.
4. "Systems Study for Control of Emissions Primary Non-ferrous Smelting
Industry, " Arthur G. McKee and Co., Vol. I, Section VIII, Contract
PH 86-65-85.
5. "Conceptual Design and Cost Study Sulfur Oxide Removal from Power
Plant Stack Gas, Magnesia Scrubbing, Regeneration: Production of
Concentrated Sulfuric Acid," Tennessee Valley Authority, prepared
for EPA, PB-22509.
293
-------
Table C-l. AUXILIARY SULFURIC ACID PLANT UNIT USAGE AND COST DATA
(8-9% S02 in feed gas)
A. Utilities & Materials
Power: a) No gas cleaning
b) With dry gas
cleaning
Water: (cooling)
Catalyst
B. Operating Labor &
Maintenance
Labor
Maintenance
Taxes and Insurance
Basis
0.02 KWHr/lbSO?
0.025 kWH/lbSO^
0.08 gal/lbSO?
0.00002 liters/lbS02
Basis
1 man/shift
(for 365 days)
5% TCI/yr
1 2 1/2% TCI/yr
Unit Cost
?07015/KWH
$0.10/M Gal
$1.80/liter
Unit Cost
$8/hr
C. Fixed Charges 13.15% TCI/hr
Based on Capital Recovery Factor using 10% interest over 15 year life.
294
-------
AUXILIARY SULFURIC ACID PLANT
TOTAL CAPITAL INVESTMENT COSTS & TOTAL ANNUAL DIRECT OPERATING COSTS
(BASED ON 8-9% S02 TO ACID PLANT)
WITH WET GAS
CLEANING
V)
O
o
oc
LU
a.
O
O
LU
CC
ID
Z
CAPITAL COSTS
£ $1.0 Million' /
>- X
ACID SECTION WITH
DRY GAS CLEANING
AUXILIARY PLANT
ONLY.
OPERATING COST-
ACID PLANT WITH
DRY GAS CLEANING
OPERATING COSTS
OPERATING COST
AUXILIARY PLANT
ONLY
Costs: Mid 1974
10,000 LB/HR
S02 RATE (LB/HR) OF REGENERATION SYSTEM
295
FIGURE C-1
-------
Table C-2. AUXILIARY SULFURIC ACID PLANT
(No gas conditioning)
CAPITAL AND TOTAL ANNUAL COSTS
TOTAL CAPITAL INVESTMENT
Annual Cost
A. Direct Operating
1. Power
2. Water
3. Catalyst
4 . Labor
5. Maintenance
6. Taxes & Insurance
TOTAL ANNUAL DIRECT
OPERATING COST
B. Annual Capitalized
Cost
C. Net Total
Annualized Cost
$ ,
$/Annual ton
SO,, Removed
$/yr/ton
SO 2 Removed
Equivalent SO^ Control System Gas Flow Rate-SCFM @1% S02
70,000
1,420,000
56
16,500
4,400
2,000
70,100
71,000
35,500
199,500
186,700
245,000
10
100,000
1,750,000
48
23,600
6,300
2,800
70,100
87,500
43,800
234,100
230,100
295,800
8
200,000
2,675,000
37
47,100
12,600
5,700
105,100
133,800
66,900
371,200
351,800
459,200
6
300,000
3,450,000
25
70,700
18,900
8,500
105,100
172,500
86,300
462,200
453,700
583,500
5
10
\D
ON
Efficiency of Sulfuric Acid Plant 97%.
Based on Corporate Tax Rate of 48%.
-------
Table G- 3. AUXILIARY SULFURIC ACID PLANT
(Including Dry Gas Cleaning)
CAPITAL AND TOTAL ANNUAL COSTS
TOTAL CAPITAL INVESTMENT
Annual Cost
A. Direct Operating
1. Power
2. Water
3. Catalyst
4. Labor
5 . Maintenance
6. Taxes & Insurance
TOTAL ANNUAL DIRECT
OPERATING COST
B. Annual Capitalized
Cost
C. Net Total 2
Annualized Cost
$
$/ Annual ton
SO 2 Removed1
$/yr/ton
S02 Removed
Equivalent SC
70,000
2,000,000
79
19,500
4,400
2,000
70,100
100,000
50,000
246,000
263,000
326,800
13
>„ Control Systeu
100,000
2,520,000
70
27,900
6,300
2,800
70,100
126,000
63,000
296,100
331,400
404,600
11
i Gas Flow Rate
200,000
3,950,000
55
55,800
12^600
5,700
105,100
197,500
98,800
475,500
519,400
639,700
9
- SCFM @1% S02
300,000
5,400,000
50
83,700
18,900
8,500
105,100
270,000
135,000
621,200
710,100
859,500
8
to
Efficiency of Sulfuric Acid Plant 97%.
"Based on Corporate Tax Rate of 48%.
-------
APPENDIX D
AUXILIARY SULFUR PLANT
Although the Allied Chemical Corporation process for reducing SC^
to elemental sulfur will handle a wide range of SO™ concentrations, the
presence of oxygen in the gas stream increases the consumption of the
natural gas reductant and tends to distort the economics unfavorably.
Operating costs are thus minimized in situations where the S02 con-
centration approaches 100 percent. Auxiliary sulfur plants therefore
become a viable option for those regenerable S02 control processes which
have the capability of generating an SO™ stream of 80 percent concentration
or better.
Capital and annual operating cost curves based on the hourly SO™
rate of the regeneration system itself are provided in Figure D-l.
Capital costs were developed from data included in a paper presented by
William D. Hunter of Allied Chemical Corporation at the 1973 Flue Gas
Desulfurization Symposium. Operating costs were developed from usage
data provided in the report, "Applicability of Reduction to Sulfur
Techniques to the Development of New Processes for Removing SO™ from
Flue Gases," Allied Chemical Corporation, Contract PH-22-68-24. Unit
values and the applicable labor, and maintenance allowances are provided
in Table D-l.
Table D-2 has been prepared to provide a ready matching of capital
and operating costs with the primary control systems in terms of various
gas flows to the primary system at 1 percent SO™.
298
-------
Table D-l. AUXILIARY SULFUR PLANT UNIT USAGE AND COST DATA
A. Chemicals & Utilities
Basis
Unit Cost
Methane (Natural gas)
Power:
Steam credit
3.1 CF/lbS02 input
0.0096 KWHr/lbS02
0.4 lb/lbS00
$1.25/M CF
$0.015/KWH
$0.80/M Ib
B. Operating Labor &
Maintenance
Basis
Unit Cost
Labor
Maintenance
Taxes and Insurance
1 man/shift
+1 man/day
5% TCI/yr
2 1/2% TCI/yr
$8/hr
C. Fixed Charges 13.15% TCI/hr
Based on Capital Recovery Factor using 10% interest ouer 15 year life.
299
-------
AUXILIARY SULFUR PLANT
TOTAL CAPITAL INVESTMENT COSTS
& TOTAL ANNUAL DIRECT OPERATING COSTS
o
o
H-
z
LU
I
LU
E
<
O
EFFICIENCY = 92%
INPUT GAS = 85% S02
OPERATING COST
$10.0 Million
CAPITAL COST
1.0 MILLION
Cos,
10,000 LBS/HR
S02 RATE (LB/HR) OF REGENERATION SYSTEM
300
-------
Table D-2- AUXILIARY ELEMENTAL SULFUR PLANT
CAPITAL AND TOTAL ANNUAL COSTS
TOTAL CAPITAL INVESTMENT
Annual Cost
A. Direct Operating
1 . Power
2. Natural gas
(methane)
3. Steam (credit)
4 . Labor
5 . Maint enanc e
6. Taxes & Insurance
TOTAL ANNUAL DIRECT
OPERATING COST
B. Annual Capitalized
Cost
C. Net Total
Annualized Cost
$/ Annual ton
S02 Removed
$/yr/ton
S02 Removed
Equivalent S02 C
70,000
2,940,000
123
7,700
201,700
(16,000)
86,700
147,000
73,500
500,000
386,600
552,500
23
Jontrol System C
100,000
3,370,000
98
10,900
288,200
(23,700)
86,700
168,500
84,300
614,900
443,200
655,100
19
Jas Flow Rate -
200,000
4,450,000
65
21,900
576,300
(47,400)
86,700
222,500
111,300
971,300
585,200
797,100
12
- SCFM @1% S02
300,000
5,450,000
53
32,800
864,500
(71,100)
86,700
272,500
136,300
1,321,700
716,700
1,229,600
12
co
o
Efficiency of Sulfur Plant 92%.
Based on Corporate Tax Rate of 48%.
-------
APPENDIX E
SO LIQUEFACTION AND STORAGE COSTS
Capital and Annual pirect Operating Costs - Figure E-l
302
-------
LIQUEFACTION AND STORAGE OF SO2
in
te
o
o
cc
LU
Q_
O
D
Z
O
LU
tr
$100,000
$40,000
0.9
0,8
0.7
0.6
0.5
0.4
0.3
CO
DC
o
Q
UL
O
z
o
_J
0.2
A - Capital cost of liquefaction
and one day storage of
liquid sulfur dioxide
B - Direct annual operating cost
S02-lbs/hr
0.1
10,000
FIGURE E-1
100,000
303
-------
APPENDIX F
UTILITY COSTS
Water Treatment Facilities - Figure F-l
Package Boilers - Figure F-2
Electrical Substations - F-3
304
-------
$106
CAPITAL COST OF WATER TREATING FACILITIES
A • Demineralizing System
B - Softening
C - Filtering
$10.000
100
1000
WATER FLOW-Gallons Per Minute
305
FIGURE F-1
-------
CAPITAL COST PACKAGE BOILERS
100,000
CO
a:
E
to
Si
in
FIGURE
Capital Cost-Package Boilers
10,000
$150,000
$300,000
$600,000
-------
CAPITAL COST OF ELECTRICAL SUBSTATIONS
100,000
2
<
X
z
o
<
w
cc
D
10,000
1000
INSTALLED COST OF SUBSTATION
INCLUDING PAD, FENCING, SWITCHGEAR
FIGURE F-3
Costs: Mid 1974
COST-DOLLARS
X 10~
100,000
-------
APPENDIX G
Capital and Annual Operating Cost Computation Sheets for Selected S0~
Control Processes Matched to Specific Primary Copper Smelters
308
-------
Appendix G
WEAK S02 STREAM CONTROL COPPER SMELTERS
CAPITAL INVESTMENT COSTS
SMELTER; Asarco - El Paso. Texas
CONTROL PROCESS Double Alkali
BASIS; Maximum gas flow 250,000 SCFM
Average S02 rate 23yQQQ Ibs/hr
Temperature at control point 250 F
SPECIAL CONDITIONS
Roaster and reverberatorv gases are quenched and cooled in a spray chamber
prior to ESP.
COST DETERMINATION
1. Prequench Cost = 100.000
2. SO- Absorption cost (includes 5% = 2.360,000
TOTAL allowance) = 2.460,000
(a) Retrofit allowance (50% C.I.) = 1.230,000
TOTAL =" 3,690,000
3. S02 Handling Section cost (21,8001b/ht) 6,700,000
(a) Disposal =
TOTAL ~ 6,700,000
4. Auxiliary plant
(a) Liquid S02 = .
5. Support Services
(a) Power (3.300KVA) - 350,000
(b) Steam ( Ibs/hr) -
(b) H2S04 • N/A
(c) Modifications to existing »
H2SO, plant
TOTAL BOUNDARY LIMIT COSTS - 10,390.000
(c) Water (0.2 Mgal/min) - 200.000
TOTAL - 550.000
6. Special costs -
• (a) ESP "
(b) Alternate processing equip. -
(c) General site costs @20% Item 4 . 2,080,000
TOTAL - 2»080>OOP
TOTAL CAPITAL INVESTMENT - $13.020,OOQ_
309
-------
WEAK SO. STREAM CONTROL COPPER SMELTERS
ANNUAL OPERATING COSTS
Smelter; Asarco-El Paso, Texas
Basis:
Control Process:
Appendix G
Double Alkali
Max. Gas Flow
Av. S02 Rate
COST COMPUTATION;
1. Gas Conditioning & S02
Absorption
a. Power
b. Make-up water
c. Neutral. Limestone
TOTAL
2. S02 Handling (21,800 Ib/hr)
a. Chemicals
b. Power
c. Steam
d. Fuel Oil or Nat.. Gas
e. Water
f. Disposal
TOTAL
3. Labor (2 1/4 man/shift)
4. Maintenance
a. Gas Condn. @ % TCI
b. S02 Handling @ 4 % TCI
5. Insurance & Taxes @ 2 1/2% TCI
250,000 SCFM
Temp, at Control Point_
250°F
23.000 Ibs/hr
Basis
i.O KWh/MSCFM
L.Ogal/MSCFM
3.021b/MSCFM
See appropria
(Uii5jafliyihRO.
19710 hr
$JLO,390,000
$13,020,000
Unit Cost
$ 0.015/KWh
$ 0.3/Mgal
$ 8/ton
:e process sectio
L6 0.015/KWh
-1.0.3/Mgal
_.$__3/ton
$ 8/hr^
Annual Cost
122,000
37,000
20,000
179,000 _^
n 2,753,000
200,000
____942L000 ,,
3.906T000 ^^
158,000 >
416,000
326,000
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
$ 4.985,000
a.
b. Alternative
N/A
c. Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost
8. Total Net Annualized Cost
9. Annual Cost/ton SO. Removed
$ 4.985,000
1,712,000
3.888,000
$44
310
-------
Appendix G
WEAK S02 STREAM CONTROL COPPER SMELTERS
CAPITAL INVESTMENT COSTS
SMELTER: Asarco - El Paso. Texas
CONTROL PROCESS Magnesium Oxide
BASIS; Maximum gas flow 250»000 SCFM
Average SO™ rate 23,000 " ' ' Ibs/hr
Temperature at control point
SPECIAL CONDITIONS
250 °T
Roaster and reverberatory gases are quenched and cooled in a spray chamber
prior to ESP.
COST DETERMINATION
1. Prequench cost = 100,000
2. S02 Absorption cost (includes 5% = 2.680rOOP
TOTAL allowance) = 2.780.OOP
(a) Retrofit allowance (50% C.I.) = 1.390.OOP
TOTAL ' = 4.17P.POP
3. S02 Handling Section cost(20,7001b/hi^
(a) Disposal =
TOTAL = 5.10Q.OPO
Auxiliary plant
(a) Liquid S02 =
(b) H2S04 - 4f350.OOP
(c) Modifications to existing =
H_SO, plant
TOTAL BOUNDARY LIMIT COSTS =* 13.62P.PPQ
5. Support Services
(a) Power (4000_KVA) - 370.000
(b) Steam ( Ibs/hr) =
(c) Water (0.2 Mgal/roin) = 200.000
TOTAL " 57P.OPP
6. Special costs —
' (a) ESP "
(b) Alternate processing equip. •
(c) General site costs @20% item A - 2.720.POO
TOTAL " 2.720.000
TOTAL CAPITAL INVESTMENT « $16,910,000
311
-------
WEAK S02 STREAM CONTROL COPPER SMELTERS
n _ F.I Paso
Smelter; ^
Basis: Max. Gas Flow
ANNUAL OPERATING COSTS
Control Process:
Appendix G
Double Alkali
120.000 SCFM
Temp, at Control Point
250°F
Av. SO- Rate_
13.800
COST COMPUTATION;
1. Gas Conditioning & S02
Absorption
a. Power
b. Make-up water
c. Neutral. Limestone
TOTAL
2. S02 Handling (13,100 Ib/hr)
a. Chemicals
b. Power
c. Steam
d. Fuel Oil or Nat. Gas
e. Water
£. Disposal
TOTAL
3. Labor (2 1/4 man/shift)
4. Maintenance
a. Gas Condn. @ % TCI
b. SO. Handling @ 4 % TCI
/
5. Insurance & Taxes @ 2 1/2% TCI
Basis
4.0KWh/MSCFM
l.Ogal/MSCFM
0.021b/MSCFM
5ee appropria
0.075KWh/lbS(
^3 S^lh/sn
1
-LSL^ZlUJirs—
A? 9«n nnn
$9,160,000
Unit Cost
$ 0.015/KWh
$ 0.3/Mgal
$ 8/ton
:e process sectic
I $ 0.015/KWh
-$_-37±OJX._ _ _.
__J_.8_/Jxr_
Annual Cost
59,000
18,000
5,000
82,000
n 1,654,000
120,000
_ _ 5_&6X000 __ _
2.340,000
158 .,.000
291^000
229,000
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
$ 3,100,000
a.
b. Alternative
N/A
c. Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost
8. Total Net Annualized Cost
$ 3,100,000
9. Annual Cost/ton SO. Removed
$ 1,205,000
$ 2.524.000
$4,7
312
-------
Appendix G
WEAK S02 STREAM CONTROL COPPER SMELTERS •
CAPITAL INVESTMENT COSTS
SMELTER; Asarco - El Paso. Texas
CONTROL PROCESS Magnesium Oxide
BASIS; Maximum gas flow 120,000 SCFM
Average S02 rate 13; 800' Ibs/hr
Temperature
SPECIAL CONDITIONS
Temperature at control point 250 °F
Treating reverberatory gases only. (Roaster gases mixed with lead sintering
machine goes to acid plant - planned)
COST DETERMINATION
1. Crequench cost = 70,000
2. SO, Absorption cost (includes 5% = 1,760,000
2 allowance)
TOTAL = i
(a) Retrofit allowance (50% C.I.) - 920,000 _
' TOTAL - 2,750,000
3. S02 Handling Section cost U2,4001b/hE) 3.600.000
(a) Disposal = _
TOTAL = 3,600.000
.,4 . Auxiliary plant
(a) Liquid S02 " _
(b) H2S04 - 3>100'000
(c) Modifications to existing = _
. H2SO, plant
TOTAL BOUNDARY LIMIT COSTS - 9,450,000
5. Support Services
(a) Power (220° KVA) = 300'00°
(b) Steam ( Ibs/hr) . = :
(c) Water ( Q.I Mgal/roin) - 120.000
TOTAL " 420.000
6. Special costs ~~ -
(a) ESP * !
(b) Alternate processing equip. ™ .
(c) General site costs @20% item 4 - 1.890.000
TOTAL " 1.890.OOP
TOTAL CAPITAL INVESTMENT - $11.760.000
313
-------
Appendix G
WEAK SO- STREAM CONTROL COPPER SMELTERS
ANNUAL OPERATING COSTS
Smelter: Asarco - El Paso, Texas
Basis: Max. Gas Flow
Control Process; Magnesium Oxide
Av. S0_ Rate
250.000 SCFM Temp, at Control Point_
23.000 Ib/hr
250°F
COST COMPUTATION;
1. Gas Conditioning & SO.
Absorption
a. Power
b. Make-up water
c. Neutral. Limestone
TOTAL
2. SO. Handling (20,700 Ib/hr)
a. Chemicals
b. Power
c. Steam
d. Fuel Oil or Nat. Gas
e. Water
f. Disposal
TOTAL
3. Labor ('3 1/2 man/shift + 1 day)
4. Maintenance
a. Gas Condn. @ Z TCI
b. S00 Handling @ 5 % TCI
£. -'"
5. Insurance & Taxes @ 2 1/2% TCI
Basis
4.0KWh/MSCFM
l.Ogal/MSCFM
0.021b/MSCFM
See appropri<
0.1KWh/lbS02
0 fm a1/1b$<
-£L_2eal/lbSC;
^
3_2j_740 hrs
$ 9.270,000
$12,560,000
Unit Cost
$ 0.015/KWh
$ 0 . 3/Mgal
$ 8/ton
te process
$ 0.015/KWh
I £ 0 3&al "
f a O'. 3/Mgal
.§„ 8/hr
Annual Cost
122,000
37,000
20,000
179,000 _
404,000
253,000
1., 8 7 5^000
lOjOOO
2,542,000
262,000 _^
464,000
314,000
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
^SO, Plant
3,761,000
a.
520,000
b. Alternative
c
Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost
8. Total Net Annualized Cost
9. Annual Cost/ton SO- Removed
314
$4.281,000
$2,224,000
$3.909,000
$48
-------
Appendix G
WEAK S02 STREAM CONTROL COPPER SMELTERS
CAPITAL INVESTMENT COSTS
SMELTER! Asarco- El Paso, Texas
CONTROL PROCESS Citrate
BASIS; Maximum gas flow 250tOOO SCFM
Average S02 rate 23^000 Ibs/hr
Temperature at control point 250 F
SPECIAL CONDITIONS
Roaster and reverberatory gases are quenched and cooled in a spray chamber
prior to ESP.
COST DETERMINATION
1. Prequench Cost = 100,OOP
2. SO. Absorption cost(includes 5% = 2,360,000
allowance) „ lrn ___
TOTAL = 2,460,000
(a) Retrofit allowance(50% C.I.) - '
TOTAL = - 3,690,000
3. SO- Handling Section cost(21,8001b/h^)
(a) Disposal
TOTAL - 6,700,000
4. Auxiliary plant
(a) Liquid S02
5. Support Services
(a) Power (3300 KVA) = 350.000
(b) Steam ( Ibs/hr) - .__
(c) Water (0-8 Mgal/min) = 450.000
6. . Special costs
(a) ESP
(b) Alternate processing equip.
(c) General site cosfcs@20Z item 4 » ?infln]nnn
TOTAL "
(b) H2S04 - N/A
(c) Modifications to existing •
H2SO, plant
TOTAL BOUNDARY LIMIT COSTS = 10.390,000
TOTAL = _ «nnnnn
TOTAL CAPITAL INVESTMENT - $13r27QfQQQ
315
-------
WEAK S02 STREAM CONTROL COPPER SMELTERS
ANNUAL OPERATING COSTS
Smelter; Asarco - El Paso, Texas
Basis: Max. Gas Flow 250,OOQSCFMTemp. at Control polnt 250°F
Appendix G
Control Process: Citrate
Av. S02 Rate_
COST COMPUTATION:
23,000 Ib/hr
1. Gas Conditioning & SO-
Absorption
a. Power
b. Make-up water
c. Neutral. Limestone
TOTAL
2. S02 Handling (21,800 Ib/hr)
a. Chemicals
b. Power
c. Steam
d. .Fuel Oil or Nat. Gas
e. Water (a) Process
(b) Cooling
TOTAL
3. Labor (4 man/shift)
4. Maintenance
a. Gas Condn. @ % TCI
b. S02 Handling @_4_% TCI
5. Insurance & Taxes @ 2 1/2% TCI
Basis
4.0KWh/MSCFM
Unit Cost | Annual Cost
$ 0.015/KWh I 122,000
Jipop.
20,000
179,000
See appropriate process
-sectioj _l_,_0_5_6_,_00j)
JkQlS^^bsL $0^015/KWh 1 _20_0_,_0_00
l. 25/MCF801,000
80,~000~
0.1/Mgal 89~~o"o~0~
35,040 hr
2,226,000
$ 8/ton 280,000
10^390,000
13,270,000
416,000
332,000
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
3,433,000
a.
Alternative
c. Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost
8. Total Net Annualized Cost
9. Annual Cost/ton SO. Removed
* 316
N/A
3,433,000
1,745,000
3,106,000
$35
-------
Appendix G
WEAK S02 STREAM CONTROL COPPER SMELTERS
CAPITAL INVESTMENT COSTS
SMELTER: Asarco - El Paso, Texas
CONTROL PROCESS Double Alkali
BASIS; Maximum gas flow 120.000 SCFM
Average S02 rate 13 j800 Ibs/hr
Temperature at control point 250 F .
SPECIAL CONDITIONS
Treating reverberatory gases only. (Roaster gases mixed with lead sintering machine
goes to acid plant - planned).
COST DETERMINATION
1. Prequench cost = 70.000
2. S0_ Absorption cost (includes 5% = 1,580,000
allowance) , ,,._ nnn
TOTAL = 1,650,000
(a) Retrofit allowance (50% C.I.) = 830,000
TOTAL " = 2,480,000
3. S02 Handling Section cost (13,100 lb/hr 4,800.000
(a) Disposal = —
TOTAL = 4,800,000
4. Auxiliary plant
(a) Liquid S02 "
(b) H2S04 • *l±
(c) Modifications to existing =
H.SO, plant
TOTAL BOUNDARY LIMIT COSTS - 7,280,000
5. Support Services
(a) Power (2.000KVA) - 300»000
(b) Steam ( Ibs/hr) . =
(c) General site costs @20% item 4 . 1.460,000
(c) Water (0.1 Mgal/min) = 120'000
TOTAL = 420'000
6. Special costs —
(a) ESP "
(b) Alternate processing equip. •
TOTAL " 1,460.000
TOTAL CAPITAL INVESTMENT - $ 9,160,000
317
-------
Appendix G
WEAK S02 STREAM CONTROL COPPER SMELTERS
ANNUAL OPERATING COSTS
Smelter: Asarco - El Paso. Texas
Control Process; Magnesium Oxide
Basis: Max. Gas Flow
Av. SO. Rate_
120,000 SCFM Temp, at Control Point_
13,800 Ib/hr
250°F
COST COMPUTATION;
1. Gas Conditioning & S0_
Absorption
a. Power
b. Make-up water
c. Neutral. Limestone
TOTAL
2. S02 Handling (12,400 Ib/hr)
a. Chemicals
b. Power
c. Steam
d. Fuel.Oil or Nat. Gas
e. Water
f. Disposal
TOTAL
3. Labor (3 1/2 man/shift + 1 day)
4. Maintenance
a. Gas Condn. @
TCI
b. SO. Handling [SfTM
See approprii
O.lKWh^lfeSO
2
0.037gal/lbS(
0.2gal/lbSO.,
X'
32,740 hr
$ 6,350,000
$ 8,660,000
Unit Cost
$ 0.015/KWh
$ Qt VMgfO
S &/ ton
te process
$ 0.015/KWh
§ 0.3/ial
^ 0.3/Mg_al
$ 8/hr
Annual Cost
$ 59,000
18 QQQ
5 QQQ
82,000 _
242^000
152^000
6,000
1,523,000 ^
260,000 ^
318,000
217,000
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
a. Acid Plant
b. Alternative
2.400,000
370,000
c. Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost
8. Total Net Annualized Cost
$ 2,770,000
9. Annual Cost/ton SO- Removed
$ 1,546,000
$ 2,611,000
$52
318
-------
Appendix G
WEAK S02 STREAM CONTROL COPPER SMELTERS
CAPITAL INVESTMENT COSTS
SMELTER! Asarco - El Paso, Texas
CONTROL PROCESS Citrate
BASIS; Maximum gas flow 120.000 SCFM
Average S02 rate 13,:8QO:' Ibs/hr
Temperature
SPECIAL CONDITIONS
Temperature at control point 250 F
Treating reverberatory gases only. (Roaster gases mixed with lead sintering
machine goes to acid plant - planned )
COST DETERMINATION
1. Prequench cost = 70,000
2. SO, Absorption cost(includes 5% = 1,580.000
allowance) , ,en ___
TOTAL = 1.650.000
(a) Retrofit allowance (50% C.I.) = 830,000
TOTAL - = 2.480.000
3. SO Handling Section cost(13,1001b/hi^ 5.000.000
(a) Disposal = 100.000 '•
TOTAL = 5,100.000
4. Auxiliary plant
(a) Liquid S02 "
(b) H2SOA -
(c) Modifications to existing =
H2SO, plant
TOTAL BOUNDARY LIMIT COSTS = 7.580.000
5. Support Services
(a) Power (2000KVA) - 300'00°
(b) Steam ( Ibs/hr) =
(c) Water ( 0.* Mgal/min) = 300'000
6. Special costs
(a) ESP
(b) Alternate processing equip
TOTAL
60°>000
(c) General site costs@20% item 4 . 1.520.000
1,520,000
TOTAL CAPITAL INVESTMENT " $9.700,000_
319
-------
WEAK SO. STREAM CONTROL COPPER SMELTERS
Appendix G
ANNUAL OPERATING COSTS
Asarco - El Paso, Texas Control Process: Citrate
Smelter:
Basis: Max. Gas Flow
120.000 SCFM
Temp, at Control Point
250°F
Av. SO- Rate
13,800 Ib/hr
COST COMPUTATION;
1. Gas Conditioning & S0_
Absorption
a. Power
b. Make-up water
c. Neutral. Limestone
TOTAL
2. S02 Handling (13,100 lb/br)
a. Chemicals
b. Power
c. Steam
. d. Fuel Oil or Nat. Gas
e. Water Process
f. Disposal Cooling
TOTAL
3. Labor (4 man/shift)
4. Maintenance
a. Gas Condn. @ Z TCI
b. S02 Handling @ 4 % TCI
5. Insurance & Taxes @ 2 1/2% TCI
Basis
4.0KWh/MSCFM
See apprpria
3.6CF/lbSO~
—__._.
-1. • _Jy^d-L/ J.UkJ\Jn
$7^580^000
$9,700,000
Unit Cost
$ 0.015/KWh
S__Q
L_8ltQja.
e process
ction
.JLljJS/MCF
_i_0_:.3/Mgal
$ 8/hr
Annual Cost
59,000
82,000
_6_3_5_ippp_
48,_000_ -
53f_000 -—
1,337,000-
280,000 ,-
303^00
243,000
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
a.
2,245,000
b. Alternative
N/A
c. Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost
8. Total Net Annualized Cost
9. Annual Cost/ton S0_ Removed
$2.245.001
$1,276.000.
$2,133,000.
$40
320
-------
Appendix G
WEAK S02 STREAM CONTROL COPPER SMELTERS
CAPITAL INVESTMENT COSTS
SMELTER: ASARCO - Hayden, Arizona
CONTROL PROCESS - Double Alkali
BASIS; Maximum gas flow _ 400,000 SCFM
Average S02 rate _ 26»50Q.i :. Ibs/hr
Temperature at control point ' _ °F
SPECIAL CONDITIONS
Roaster and reverberatory gases are mixed prior to ESP. As a result of spray cooling
of the reverberatory gases, temperatures are moderate and full gas conditioning treatment
is not required.
COST DETERMINATION
1. Prequench Cost = 100.000 _
5. Support Services
(a) Power (4.300KVA) - 380,000
(b) Steam ( Ibs/hr) - .
2. SO- Absorption cost(includes 5 % = 2.990.000
TOTAL allowance) m 3to9Qio0o
(a) Retrofit allowance (50% C.I.) = 1,550,000
TOTAL " " 4,640,000
3. S02 Handling Section co(s£'°°01b/hr) - 7,300.000
(a) Disposal = .
TOTAL = 7,300,000
4. Auxiliary plant
(a) Liquid SO., •
(b) H2S04 " N/A
(c) Modifications to existing •
H-SO, plant
TOTAL BOUNDARY LIMIT COSTS - 11,940,000
(c) Water (0.2 Mgal/min) = 20Q»000
TOTAL - 580,000
6. Special costs — _
(a) ESP - '
(b) Alternate processing equip. -
(c) General site costs -@20% Item 4 * 2.390.OOP
TOTAL " 2,3Qn,nnn
$ 14,910,000
TOTAL CAPITAL INVESTMENT
321
-------
WEAK SO. STREAM CONTROL COPPER SMELTERS
ANNUAL OPERATING COSTS
Smelter; ASARCO-Hayden. Arizona Control Process:
Basis: Max. Gas Flow 400.000 SCFM Temp> afc Control Point
i£ cnn 11~ /V...
Av. S02 Rate
COST COMPUTATION:
Appendix G
Double Alkali
270°F
26,500 Ib/hr
1. Gas Conditioning & S02
Absorption
a. Power
b. Make-up water
c. Neutral. Limestone
TOTAL
2. S02 Handling (25, 000 Ib/hr)
a. Chemicals
b . Power
c. Steam
d. Fuel Oil or Nat. Gas
e. Water
f. Disposal
TOTAL
3. Labor (2 1/4 men/shift)
4. Maintenance
a. Gas Condn.
% TCI
b. S02 Handling T? _ % TCI
5. Insurance & Taxes @ 2 1/2% TCI
Basis
4.0KWh/MSCFM
l.Ogal/MSCFM
0.021b/MSCFM
See appropria
0>075iMbT
-
12.xIlQ.Jirs.
$11.940.000
$14,910,000
Unit Cost
$0.015/KWh
$0.3/Mgal
$8/ton
:e process sectio
$0.015/KWh
$0 . 3/Mgal
^3^ ton
Annual Cost
196,000
59,000
15,000
270,000 _
\ 3,157,000
230,000
-
-
12^000
1 JDSO.JJOO
4,479,000 __
_____15_8_lOp_0
478J300
373,000 ^^^
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
5,758,000
a.
b. Alternative
N/A
c. Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost
8. Total Net Annualized Cost
9. Annual Cost/ton SO- Removed
* 322
$5,758', 000
$1.961.000
4,478,000
-------
WgAJL_SQo STREAM CONTROL COPPER SMELTERS Appendix G
CAPITAL INVESTMENT COSTS
SMELTER; ASARCO - Hayden, Arizona
CONTROL PROCESS Magnesium Oxide
BASIS; Maximum gas flow _ 400.000 _
Average S02 rate _ 26. 500 i :.i Ibs/hr
Temperature at control point 270 °p
SPECIAL CONDITIONS
Roaster and reverberatory gases are mixed prior to ESP. As a result of spray
cooling of the reverberatory gases, temperatures are moderate.
COST DETERMINATION
1. Prequench Cost = _ 100,000
2. SO. Absorption cost(includes 5% = 3,410,000
TOTAL allowance) = 3f510fOQD
4. Auxiliary plant
(a) Liquid S02
6. Special costs -
(a) ESP
(b) Alternate processing equip.
(a) Retrofit allowance (25% C.I.) = 880,000
TOTAL - ' 4,390,000
(23,800 lb/hr; —
3. S0_ Handling Section cost = 5,600,000
(a) Disposal =
TOTAL = 5,600,000
(b) H2S04 = 4,750,000
(c) Modifications to existing =
HSO plant ' -
TOTAL BOUNDARY LIMIT COSTS - 14.740,000
5. Support Services
(a) Power (5000 KVA) - _ 410.000
(b) Steam ( _ Ibs/hr).
(c) Water (0.2 Mgal/min) - _ 200,000
TOTAL = 610,000
(c) General site costs @20% Item 4 - 2,950,000
TOTAL - 2,950,000
TOTAL CAPITAL INVESTMENT - $18,300,000
323
-------
WEAK SO- STREAM CONTROL COPPER SMELTERS
ANNUAL OPERATING COSTS
Smelter: ASARCO-Hayden, Arizona
Basis: Max. Gas Flow
Control Process:
Appendix G
Oxide
400,000 SCFM
Temp, at Control Point
270°F
Av. S0_ Rate_
26,5000 Ib/hr
COST COMPUTATION;
1. Gas Conditioning & SO.
Absorption
a. Power
b. Make-up water
c. Neutral. Limestone
TOTAL
2. S02 Handling (23,800 Ib/hr)
a. Chemicals
b. Power
c. Steam
d. Fuel Oil or Nat. Gas
e. Water
f. Disposal
TOTAL
3. Labor (3 1/2 man/shift*+ 1 day)
4. Maintenance
a. Gas Condn. @ - % TCI
b. S02 Handling @ - % TCI
5. Insurance & Taxes @ 2 1/2% TCI
Basis
4.0KWh/MSCFM
l^Ogal/MSCFM
0.021b/MSCFM
5ee approprial
).!KWh/lbS02
D~037gal/lbSO-
O.JZgal/lbSO..
-• " • ' •£•"
32,740 hrs
$ 9,990,000
$13,550,000
Unit Cost
$0.015/KWh
__$.0._3_/M__gal_
$8 /ion
e process sectio
$0.015/KWh
$0.30/gal
$0.3/Mgal
$8/hr
Annual Cost
196,000
59_,000_
15,000
270,000
i 464,000
291,000
_
2,156,000
12,000
—
2,923,000
262,000
500,000
339,000
^
^
^
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
H2S04 Plant
$ 4,294,000
570,000
b. Alternative
c. Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost (basis $18,300,000)
8. Total Net Annualized Cost
9. Annual Cost/ton S0_ Removed
$ 4,864,000
$ 2,407,000.
$ 4,350,000
$46
324
-------
Appendix G
WEAK SO,, STREAM CONTROL COPPER SMELTERS •
CAPITAL INVESTMENT COSTS
SMELTER: ASARCO - Hayden, Arizona
CONTROL PROCESS Citrate
BASIS: Maximum gas flow 400.000 SCFM
Average S02 rate 26.5Q.Q ;, lbs/hr
Temperature at control point ^70 °p
SPECIAL CONDITIONS
Roaster and reverberatory gases are mixed prior to ESP. As a result of spray cooling
of the reverberatory gases, temperatures are moderate and full gas conditioning treatment
is not required.
COST DETERMINATION
lf Prequench Cost ' = 100,000
2. SO. Absorption cost(includes 5% = 2,990,000
^ allowance;
TOTAL = 3,090,000
(a) Retrofit allowance (50% C.I.) = 1,550,000
5. Support Services
(a) Power ( 4300KVA) - 380,000
(b) Steam ( Ibs/hr)-. -
6. Special costs —
(a) ESP
(b) Alternate processing equip.
(c) General site costs @20% Item 4 " 2. 400 f OOP
TOTAL "
TOTAL « - 4,640,000
3. S02 Handling Section cost^25'0001b/hri 7,?nntnnn
(a) Disposal - 150,000
TOTAL = 7,350,000
4. Auxiliary plant
(a) Liquid S02
H2S°4 - T../A
(c) Modifications to existing •
HSO plant
TOTAL BOUNDARY LIMIT COSTS - 11,990,000
(c) Water ( 0.8 Mgal/rain) - 4Sn,f)OQ
TOTAL - 830,000
TOTAL CAPITAL INVESTMENT - $15.220,000
325 "~
-------
WEAK S02 STREAM CONTROL COPPER SMELTERS
ANNUAL OPERATING COSTS
Appendix G
Smelter; ASARCO-Hayden, Arizona
Basis: Max. Gas Flow 400,000 SCFM
Control Process:
Citrate
Temp, at Control Point
270°F
Av. S0_ Rate_
26,500 Ib/hr
COST COMPUTATION:
1. Gas Conditioning & S02
Absorption
a. Power
b. Make-up water
c. Neutral. Limestone
TOTAL
2. S02 Handling (25,000 Ib/hr)
a. Chemicals
b. Power
c. Steam
d. Fuel Oil or Nat. Gas
e. Water a) Process
b) Cooling
TOTAL
3. Labor (4 manshift)
4. Maintenance
a. Gas Condn. @ - % TCI
b. S0_ Handling @_4% TCI
i
5. Insurance & Taxes @ 2 1/2% TCI
Basis
4.0KWh/MSCFM
l.Ogal/MSCFM
0.021b/MSCFM
See appropria
0.075KWh/lbSO
3,6r.F/1hSOi
1.5gal/lbS02
S.Ogal/lbSO,
35,040 hrs
| ll^OJDOO
$ 15,220,000
Unit Cost
$0.015/KWh
$0.3/M gal
$8/ton
:e process sectic
, $0.015/KWh
-S-L-ZJjyMCP
$0.'3/Mgal
$0.1/Mgal
$8/hr
Annual Cost
196,000
59,000
15,000
270.000 _==»
n 734,000
230,000
qiR nnn • —
92,000
102,000
2,076,000
280,000
480JJOJ)
381,000
~***^^
^
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
3. & a;
a.
b. Alternative
c. Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost
8. Total Net Annualized Cost
9. Annual Cost/ton SO, Removed
* 326
$3,487,000
—"'
$2,001,000
$3,328,000
-------
Appendix G
WEAK S02 STREAM CONTROL COPPER SMELTERS •
CAPITAL INVESTMENT COSTS
SMELTER! ASARCO-Tacoma, Washington
CONTROL PROCESS Double Alkali
BASIS; Maximum gas flow 250.000 scpM
Average S02 rate 23.000- ' Ibs/hr
Temperature at control point 250 °F
SPECIAL CONDITIONS
Under the present system, mixed roaster and reverberatory gases are quenched and
cooled in a spray chamber before the ESP's (2 in series^.
COST DETERMINATION
1. Prequench Cost = 100,000
2. S02 Absorption cost(includes 5% " 2,lfin,nnn
TOTAL allowance) = 2,460,000
(a) Retrofit allowance (75% C.I.) - 1.840.000
TOTAL - 4,300,000
3. SO- Handling Section cost(21,800 - 6.700.000
( f _. , Ib/hr)
(a) Disposal *
5. Support Services
(a) Power ( 3300 KVA) - 350,000
(b) Steam ( Ibs/hr) »
(c) Water (0.2 Mgal/min) -
TOTAL = 6,700,000
4. Auxiliary plant
(a) Liquid SOj -
(b) H2S04 - N/A
(c) Modifications to existing =
H2SO, plant
TOTAL BOUNDARY LIMIT COSTS - J.1,000,000
TOTAL - 550»OQ0
Special costs
(a) ESP - -
'(b) *Alternate processing equip. - 6nntnnn
(c) General site costs @20% Item 4 - 2^2001OOP
TOTAL - 2,800,QQQ
TOTAL CAPITAL INVESTMENT - $14,350,000
*Bag filter (§40,000 SCFM. 327 "~~~~
-------
WEAK S02 STREAM CONTROL COPPER SMELTERS
ANNUAL OPERATING COSTS
Smelter; ASARCO-Tacoma , Washington
Basis:
Control Process:
Appendix G
Double Alkali
Max. Gas Flow
Av. S02 Rate
COST COMPUTATION;
1. Gas Conditioning & S02
Absorption
a. Power
b. Make-up water
c. Neutral. Limestone
TOTAL
2. S02 Handling (21,800 Ib/hr)
a. Chemicals
b. Power
c. Steam
d. Fuel Oil or Nat. Gas
e. Water
f. Disposal
TOTAL
3. Labor (2 1/4 man/jshift)
4. Maintenance
a. Gas Condn. @ - % TCI
b. S02 Handling @ 4 % TCI
5. Insurance & Taxes @ 2 1/2% TCI
250,000 SCFM
Temp, at Control Point
250°F
23,000 Ib/hr
Basis
0.21b/MSCFM
See appropria
te process section
0,057
KWh/lbSO
_.jmb/lbS02_
4.JLLJ1QILDJDD.
$ 14,350,000
Unit Cost
$0.015/KWh
$8/ton
$0.015/KWh
$0;3/Mgal
$3/ton.
$8/hr
Annual Cost
1221000
37_,000
20,000
179,000
200,000
11 ,_000 ^
____9_4_2_,_00_p ,-
3.906.000
158,000
359,000
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
a.
b. Alternative
JJ/A.
c. Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost
8. Total Net Annualized Cost
9. Annual Cost/ton S0_ Removed
328
-------
Appendix G
WEAK SO,, STREAM CONTROL COPPER SMELTERS •
CAPITAL INVESTMENT COSTS
SMELTER: ASARCO - Tacoma, Washington
CONTROL PROCESS Magnesium Oxide
BASIS; Maximum gas flow 250,000
Average S02 rate 23.000:. • :.. lbs/hr
Temperature at control point 250 °j?
SPECIAL CONDITIONS
Under the present system,mixed roaster and reverberatory gases are quenched and
cooled in a spray chamber bsfore the ESP's (2 in series).
COST DETERMINATION
1. Prequench Cost m 100,000
ss
(includes 5%
nnn
.000
2. S02 Absorption costanowance)
TOTAL . - 2.780.000
(a) Retrofit allowance (75% C.I.) « 2.080.000 _
TOTAL - 4,860.000
3. S02 Handling Section cost(20»7001b/h^ 5.100.000 _
(a) Disposal • _ _____
TOTAL - 5,100.000
4. Auxiliary plant
(a) Liquid S02
(b) H2S04 - 4,350,000
(c) Modifications to existing »
H2SO^ plant
TOTAL BOUNDARY LIMIT COSTS - 14,310,000
5. Support Services
(a) Power (4000 KVA) - 370'000
(b) Steam ( lbs/hr). -
(c) Water ( 0.2 Mgal/min) - 200»000
TOTAL - 570'000
6. Special costs _
(a) ESP - j
'(b) Alternate processing equip. - 600.000
(c) General site costs@20% Item 4 * 2.860.000
TOTAL - 3.460.000
TOTAL CAPITAL INVESTMENT - $18.340,000
329 . =
-------
WEAK S02 STREAM CONTROL COPPER SMELTERS
Appendix G
•\
ANNUAL OPERATING COSTS
Smelter: ASARCO-Tacoma, Washington ' Control Process: Magnesium Oxide
Basis:
Max. Gas Flow
Av. SO- Rate_
250,000 SCFM
Temp, at Control Point_
250°F
23,0001b/hr
COST COMPUTATION;
1. Gas Conditioning & SO.
Absorption
a. Power
b. Make-up water
c. Neutral. Limestone
TOTAL
2. S02 Handling (20,700 Ib/hr)
a. Chemicals
b. Power
c. Steam
d. Fuel Oil or Nat. Gas
e. Water
f. Disposal
TOTAL
3. Labor (3.S/2 man/sbift + 1 day)
4. Maintenance
a. Gas Condn. @ ~ % TCI
b. S02 Handling @ 5 % TCI
5. Insurance & Taxes @ 2 1/2% TCI
Basis
4.0KWh/MSCFM
l.Ogal/MSCFM
0.021b/MSCFM
See appropriat
0.1KWh/lbS02
0.2_ga_l/lbSp5J
$ 32,740 hrs
$ 9JJM^M_
$ 13,990,000
Unit Cost
$0.015/KWh
$0.3/Mgal
$8/ton
e process sectioi
. $0.015/KWh
icf a/ ai
_ 10 '. 3/M _gal .
$8/hr
Annual Cost
122,000
37,000
20,000
179,000 _^,
404,000
253,000
i SZ5. aaa
10,000
2,542,000
262,000
y^QQQ
350,000
— -
, "
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
a. H?SOA Plant .
520,000
b. Alternative
c. Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST-.
7. Annualized Capital Cost (basis $18,340,000)
8. Total Net Annualized Cost
$ 4.351iOQ
$ 2,412,000
9. Annual Cost/ton S0_ Removed
330
-------
Appendix G
WEAK S02 STREAM CONTROL COPPER SMELTERS •
CAPITAL INVESTMENT COSTS
SMELTER; ASARCO-Tacoma, Washington
CONTROL PROCESS Citrate
BASIS; Maximum gas flow 250,000 SCFM
Average S02 rate 23,000 :..! :,. lbs/hr
Temperature at control point 250 °p
SPECIAL CONDITIONS
Under the present system, mixed roaster and reverberatory gases are quenched and
cooled in a spray chamber before the ESP's (2 in series)
COST DETERMINATION
1. Prequench Cost .. = 100,000
_ _„ ., (includes 5% 0 „,.
2. S02 Absorption cost allowance) - 2'360'
TOTAL
(a) Retrofit allowance (75% C.I.) • 1 .RAO. OOP _
TOTAL = 4,300,000
(21,800 Ib/hr) ~~ ~
3. S0_ Handling Section cost = 6(7QQ,QQp _
(a) Disposal = _ _
TOTAL = 6.700.000
4. Auxiliary plant
(a) Liquid S02 - _
(b) H2S04 = N/A
(c) Modifications to existing =
HSO plant
TOTAL BOUNDARY LIMIT COSTS = 11,000,000
5. Support Services
(a) Power (3300 KVA) -
(b) Steam ( lbs/hr)
(c) Water (°^_Mgal/min) - 650,000
TOTAL - 800.000
6. Special costs _
(a) ESP --
.(b) Alternate processing equip. - 600'000
(c) General site costs@20% Item 4 • 2r200rOOP
TOTAL - $ 2,800,000
TOTAL CAPITAL INVESTMENT - $ 14,600,000
331 =
-------
WEAK S02 STREAM CONTROL COPPER SMELTERS
ANNUAL OPERATING COSTS
Smelter; ASARCO-Tacoma. Washington ' Control Process:
Basis: Max. Gas Flow 250.000 SCFM
Appendix G
Citrate
Temp, at Control Point__250F
Av. SO. Rate_
23,000 Ib/hr
COST COMPUTATION;
1. Gas Conditioning & S0_
Absorption
a. Power
b. Make-up water
c. Neutral. Limestone
TOTAL
2. S02 Handling (21,800 Ib/hr)
a. Chemicals
b. Power
c. Steam
d. Fuel Oil or Nat. Gas
e. Water a) Process
b) Cooling
TOTAL
3. Labor (4 man/shift)a
4. Maintenance
a. Gas Condn. @_J^_% TCI
b. S0£ Handling @ 4 % TCI
5. Insurance & Taxes @ 2 1/2% TCI
Basis
i.OKWh/MSCFM
L021b/MSCFM_
See ap_propria
0.075KWh/lbSO
S.eCF/lbSO^
1.5 gal/lbS02
5.0 gal/lbSO,
$ 35^040 hrs
?11 ,000,00"
514,600,000
Unit Cost
$0.015/KWh
_JP_._3/Mga_l_
__$_8 /ton
e process sectio
-$0.015/KWh
"$1.25/MCF
$0/3/M gal
$0.1/M.gal
$8/hr
Annual Cost
122,000
20,00p_
179.000 _=*•
i 1,056,000 --
200,000 ___.-
__^
801,000 -
80,000
89,000 __-
2,226,000
280,000 „•
.4.4.0,000 •-'"'
365, 000^^^,^
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
3,490,000
a.
b. Alternative
_N/A_
c. Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost
8. Total Net Annualized Cost
$1,920,000
$3,268,000
9. Annual Cost/ton SO- Removed
$37
332
-------
STREAM CONTROL COPPER SMELTERS • APPendix G
CAPITAL INVESTMENT COSTS
SMELTER; Kennecott - Hayden, Arizona
CONTROL PROCESS. Double Alkali
BASIS; Maximum gas flow __ 165,000 SCFM
Average S02 rate _ :..! 3>.,000_lbs/hr
Temperature at control point _ 300 °p
SPECIAL CONDITIONS
COST DETERMINATION
1. Gas conditioning cost « 2,030,000
(a) Adjusted cost - 1,665,000
2. S02 Absorption cost (includes 5% - 1.890.000
TOTAL allowance) „ 3.555.OOP
(a) Retrofit allowance (20%C.I.) - 710.000
TOTAL = " 4.270.000
3. S02 Handling Section cost (2850 Ib/hr^ 1.850.000
(a) Disposal "
TOTAL * 1.850.000
4. Auxiliary plant ,
(a) Liquid S02
H2S04 " N/A
(c) Modifications to existing «
H2SO^ plant
TOTAL BOUNDARY LIMIT COSTS - 6.120.000
5. Support Services
(a) Power ( 2200 KVA) - 295.000
(b) Steam ( Ibs/hr). -
(c) Water (0.6 Mgal/min) - 380.000
TOTAL ' • 675,000
6. Special costs —
(a) ESP - -
(b) Alternate processing equip. »
(c) General site costs @20% item 4 - 1,224,000
TOTAL - 1,224.000
TOTAL CAPITAL INVESTMENT - 8.020,000
333 ~~
-------
WEAK S02 STREAM CONTROL COPPER SMELTERS
Appendix G
ANNUAL OPERATING COSTS
Smelter; Kennecott-Hayden, Arizona * Control Process: Double Alkali
Basis:
Max. Gas Flow 165.000 SCFM
Temp, at Control Point 30° F
Av. SO, Rate_
3,000 Ib/hr
COST COMPUTATION:
1. Gas Conditioning & SO,
Absorption
a. Power rr- ,,,-,760
[5.25
b. Make-
4.0]
up
c. Neutral. Limestone
TOTAL
2. S02 Handling (2850 Ib/hr)
a. Chemicals
b . Power
c. Steam
d. Fuel Oil or Nat. Gas
e. Water
f. Disposal
TOTAL
3. Labor (2 man/sniff) 3
A. Maintenance
a. Gas Condn. @ 5 Z TCI
b. S02 Handling @ 4Z TCI
5. Insurance & Taxes @ 2 1/2% TCI
Basis
6 . OKW/MSCFM
2 5gal/MSC!M
0.071b/MSCFM
See appropri
section
0.075KWh/lbSO
—
0.2gal/lbSO
}
17^520 hrs
£ 1 1198^000
i 4 illj^OOO
$ 8,020,000
Unit Cost
§. 0.015/KWh
$ o.30/Mgal
^ 8/ton
ite process
, $ 0,015/KWh
$ O.'3/Mgal
----J-tss-
$ 8/hr
Annual Cost
$ 121,000
61,000
23,000
205,000 _=^>.
360,000 _,-.
26,000 __^
,„
2,000
511,000
140_,000
100,000
.
165,000 _^-'
200,000 ^^ia^
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
a.
b. Alternative
c. Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized-Capital Cost
8. Total Net Annualized Cost
9. Annual Cost/ton S0_ Removed
334
-------
WEAK S02 STREAM CONTROL COPPER SMELTERS .APPendix G
CAPITAL INVESTMENT COSTS
SMELTER: Kennecott - Hayden, Arizona
CONTROL PROCESS Citrate
BASIS; Maximum gas flow 165,000 SCFM
Average S02 rate :3»OQO lbs/hr
Temperature at control point 300 p
SPECIAL CONDITIONS
COST DETERMINATION
1. Gas conditioning cost = 3.180.000
(a) Adjusted cost - 2,600.000
2. SO, Absorption cost(includes 5% - 1,890.000
TOTAL allowance> = 4.490.000
(a) Retrofit allowance (20%C.I.)
900,000
TOTAL - 5.390.000
3. S02 Handling Section cost(2850 lb/hr)= 2.250.000
(a) Disposal
20.000
6. Special costs_
(a) ESP
' (b) Alternate processing equip.
TOTAL • 2.270.000
4. Auxiliary plant
(a) Liquid S02
(b) H2S04
(c) Modifications to existing =
H-SO, plant
TOTAL BOUNDARY LIMIT COSTS = 7,660,000
5. Support Services
(a) Power (2600 KVA) - 320,000
(b) Steam ( lbs/hr).
(c) Water (1.0 Mgal/min) - 500,000
820.000
TOTAL
TOTAL
$10,010,000
TOTAL CAPITAL INVESTMENT • ===========
335
-------
WEAK S02 STREAM CONTROL COPPER SMELTERS
Appendix G
ANNUAL OPERATING COSTS
Smelter; Kennecott-Hayden, Arizona Control Process; Citrate
Basis: Max. Gas Flow 165,000 SCFM
Temp, at Control Point
300°F
Av. SO- Rate_
3,000 Ib/hr
COST COMPUTATION:
1. Gas Conditioning & S0
+ 4.0]
760 ,.(
Absorption
r, n.,760 ,3
a. PowerL7.0(I^0)
b. Make-up water [31
c. Neutral. Limestone
TOTAL
2. S02 Handling (2850 Ib/hr)
a. Chemicals
b. Power
c. Steam
. d. Fuel Oil or Nat. Gas
e. Water (a) Process
(b) Cooling
TOTAL
3. Labor (3 1/2 man/shift)
4. Maintenance
a. Gas Condn. @ 5 _% TCI
b. S0_ Handling @_j_Z TCI
5. Insurance & Taxes @ 2 1/2% TCI
Basis
6.6 KW/MSCFM
_2_-_5gal/MSC.F.M
JLPJJJb/MSCEM
See appropria
section
0.075KWh/lbSO
3.6CF/lbSO£"
^ 5 ga L/_lhSO~
5^0gal/ IbJiQn
30U660 hrs
? 4^538^000
$10,010,000
Unit Cost
§. 0.015/KWHr
$_8/tpn
e process
,$.0.015/KWh
$ 1.25/MCF
_$ _Oj_3^Mgal
$ Qil/Mgal
$ 8/hr
Annual Cost
$ 133,000
61,000
29,000
223,000
84,000
28,000
105,000
lljOOO _.
12., 000 _.
240,000
245,000
156,000
182,000
250,000
^
__„
__„
^^
^
___^-
^
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
1,296,000
a.
N/A
b. Alternative
c. Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost
8. Total Net Annualized Cost
$ 1,296,000.
$ 1,316,000_
$ 1,670,000
9". Annual Cost/ton S0_ Removed
336
-------
Appendix G
WEAK S02 STREAM CONTROL COPPER SMELTERS •
CAPITAL INVESTMENT COSTS
SMELTER; Kennecott - Hayden, Arizona
CONTROL PROCESS Citrate - Direct Stripping Option
BASIS; Maximum gas flow 165,000 SCFM
Average S02 rate 3.000 :.. Ibs/hr
Temperature at control point 300 F
SPECIAL CONDITIONS
COST DETERMINATION
1. Gas conditioning cost = 3.180.000
(a) Adjusted cost - 2.600.000
2. SO, Absorption cost (includes 5% - 1,890.000
TOTAL allowance) = 4,490.000
(a) Retrofit allowance (20% C.I.) - 900'000
TOTAL - " 5,390.000
3. SO, Stripping Section Cost = 500'000 _
(a) Disposal = -
TOTAL - _ 50Q'000
4. Auxiliary plant
- 200,000
-
(b) Steam gO.OOO.lbs/hr). - _ 13°'00°
(c) Water ( 1-0 Mgal/min) - _ 500'000
-
6. Special costs
(a) ESP
(b) Alternate processing equip.
(c) General site costs @20% item 4 - 1>220.000
/ \ T., jj on
(a) Liquid SO-
(b) H2SOA = -
(c) Modifications to existing * -
H.SO, plant
TOTAL BOUNDARY LIMIT COSTS - $6,090.000
5. Support* Services
(a) Power ( 2300KVA)
930,000
TOTAL * -
1'220'000
TOTAL CAPITAL INVESTMENT - $8>24°>00°
-------
WEAK SO, STREAM CONTROL COPPER SMELTERS
Appendix G
Smelter:_
Basis: Max. Gas Flow
Av. SO- Rate_
ANNUAL OPERATING COSTS
- Hay^nr Arizona ' Control Process: Citrate-Direct
165.000SCFM Temp, at Control Point 300 F
3,000 Ib/hr
COST COMPUTATION:
1. Gas Conditioning &
Absorption
. „, n/760 v.6-1
b. Make-up water L-HIQ^Q' J
c. Neutral. Limestone
TOTAL
2. S02 Handling (2850 Ib/hr) •
a. Chemicals
b. Power
c. Steam
d. Fuel Oil or Nat. Gas
e. Water
f. Disposal
TOTAL
3. Labor (2 4 man/shift)
4. Maintenance
a. Gas Condn. @ 5% TCI
b. SO- Handling @ * % TCI
JL TrT
5. Insurance & Taxes @ 2 1/2% TCI
Basis
6 . 6KW£MSCFM
2^5gal/MSCF_M
0.091b/MSCFM
JJILhsAbSO
- 2
_100P_JJPM__
_21j_9_gg_hr_s__
$3A120,000
^2,968,000
$8,240,000
Unit Cost
^__P_.JO/Mgal j
$ 8/ton
a _L^5/J1JJ^
_$__g'._3/_Mgal_
_$__8/hr
Annual Cost
$ 133,000
61_,_000_ -
29,000 -
223, 000__^
3,000 (Es^
29-i.aaa
i47,.ggg »—
4411000==*=^
_ 175^000 ^^>
156,000 ;—
H9aOOO —
__Jw.ooU^
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
S0_ liquifaction
b. Alternative
c. Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost
8. Total Net Annualized Cost
9. Annual Cost/ton S0_ Removed
-------
SMELTER:
WgAKjQ, STREAM CONTROL
CAPITAL INVESTMENT COSTS
Kennecott - Hurley, New Mexico
CONTROL PROCESS Double Alkali
BASIS; Maximum gas flow
Average S0_ rate
Temperature at control point
.SPECIAL CONDITIONS
270,000
20,200
_SCFM
_lbs/hr
425
ESP cannot handle present gas flows
COST DETERMINATION
1. Gas conditioning cost =*
(a) Adjusted cost «
2. S02 Absorption cost (includes 5%
TOTAL allowance)
(a) Retrofit allowance (25%C.I.)
TOTAL
3. S02 Handling Section cost (19,2001b/hF>
(a) Disposal -
TOTAL
4. Auxiliary plant
(a) Liquid S02
(b) H2S04
(c) Modifications to existing -
H2SO^ plant
TOTAL BOUNDARY LIMIT COSTS -
5. Support Services
(a) Power ( KVA)
(b) Steam ( Ibs/hr).
(c) Water ( 3.2 Mgal/min)
TOTAL
6. Special costs
(a) ESP(not chargeable to process) -
(b) Alternate processing equip. -
(c) General site costs @20% item 4 «
TOTAL
TOTAL CAPITAL INVESTMENT -
339
2.450Tpon
2.450.000
4.900,000
1.230TQQO
6.100rOOQ
N/A
Appendix G
50,000
(3
,500,000)
2
,450,000
.6,130,000
6,100.000
12,230,000
50,000
2,450,000
$14,730,000
-------
WEAK SO. STREAM CONTROL COPPER SMELTERS
^
ANNUAL "OPERATING COSTS
Appendix G
Smelter; KenneMt-f-Hnrl Py,
Basis: Max. Gas Flow 275,000 SCFM
Control Process: Double Alkali
Temp, at Control Point
Av. SO Rate_
20,200 Ibs/hr
COST COMPUTATION;
1. Gas Conditioning & SO
Absorption
Pr oc/910 X3 . . -.-.
a. Power L;5-0tjo60-' +4.0]
b. Make-up water[3 •% (Z±P_).6]
c. Neutral. Limestone
TOTAL
2. S02 Handling
a. Chemicals
b. Power
c. Steam
d. Fuel Oil or Nat. Gas
e. Water
f. Disposal
TOTAL
3. Labor (3 men/shift)
4. Maintenance
a. Gas Condn. @ 5 % TCI
b. SO- Handling @ 4 % TCI
5. Insurance & Taxes @ 2 1/2% TCI
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
a.
b. Alternative
c. Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost
8. Total Net Annualized Cost
9. Annual Cost/ton SO- Removed
Basis
7.3KW/MSCFM
2-7g^l /MSGMH
O.OTlha/MSdF'1'
See appropria
sect.
Q^7'iKWh/lhso2
Q_JJsaJL/JlhSQ~
•uow^u j£.
3.J_3J_b_nb_SQr.
26^280. Jofl._.
5 3,062,000
Li»*lfiiIlflQ._.
?14,730,000
Unit Cost
$ 0.15/KWh
$ Q 3Q/Mga.T_
$ 8/fon
te process
$ n.m s/Torh
$ o.in/Mgfli^
$ 3/ton
^^/hr_
Annual Cost
$ 246,000
_LQQ QQQ,
^8 000
393,000
2,425,000
_ _ JLJ6-^QOO.-
,9^000.
fi9Q nno
3,439,000 _=======
71 n nnn
153,000
•*A7,,OM-
368,000 ___^**
$4,930,000
$4.930.000
$1,937,000
$4,029,000
$51
340
-------
WEAK S02 STREAM CONTROL COPPER SMELTERS
CAPITAL INVESTMENT COSTS
SMELTER; Kennecott-Hurlev. New Mexico
CONTROL PROCESS. Magnesium Oxide
BASIS: Maximum gas flow ?7n,nnn SCFM
Average S02 rate 20i.2,OQ, Ibs/hr
Temperature at control point 425 F
SPECIAL CONDITIONS
ESP cannot handle present gas flows and efficiency is poor
Appendix G
COST DETERMINATION
1. Gas conditioning cost - 4.400.000
(a) Adjusted cost - 3.950.000
2. SO, Absorption cost(includes 5% - 2,780,000
TOTAL allowance) _ 6,730.000
(a) Retrofit allowance (25ZC.I.) = 1,680,000
TOTAL ' " 8,410.000
3. S02 Handling Section CostU8,2001b/h^ 4,650.000
(a) Disposal =
TOTAL - 4'65°'00°
6. Special costs
v _ (4,000,000)
(a) ESP (not chargeable to process)
4. Auxiliary planf
(a) Liquid S02 "
(b) H2S04 (324 TPD) Dry Gas cleaning, 3'900>00°
(c) Modifications to existing - -
H2SO, plant
TOTAL BOUNDARY LIMIT COSTS - 16>96°'°00
5. Support Services
(a) Power ( _ KVA) - - _ -
(b) Steam ( _ Ibs/hr) - - _
(0 water <^iM8.l,»in> - -
TOTAL " -
(b) Alternate processing equip.
(c) General site costs 20% item 4 - 3.39Q.OQQ
TOTAL - 3,390,000
TOTAL CAPITAL INVESTMENT - $20,400,000,
341
-------
WEAK S02 STREAM CONTROL COPPER SMELTERS
Appendix G
ANNUAL OPERATING COSTS
Smelter; Kennecott - Hurley.New Mexico ' Control Process: Magnesium Oxide
Basis: Max. Gas Flow 275,000 SCFM
Temp, at Control Point_
425°F
Av. S0_ Rate
20,200 Ibs/hr
COST COMPUTATION;
1. Gas Conditioning & S0_
Absorption
a. Power (X)3 + 4.0]
b. Make-up water rnw^Q N .6"!
LJX<-1060; J
c. Neutral. Limestone
TOTAL
2. S02 Handling (18,100 Ib/hr)
a. Chemicals
b. Power
c. Steam
d. Fuel Oil or Nat. Gas
e. Water
f. Disposal
TOTAL
3. Labor (4^ men/shif1? + 1 day) '
4. Maintenance
a. Gas Condn. @_5__% TCI
b. S0_ Handling @ 5 Z TCI
/ —~~
5. Insurance & Taxes @ 2 1/22 TCI
Basis
8 . 4KW/MSCFM
2.7gal/MSCFM
JUBUOKIII
See appropri
section
0.075KWh/lbSO
0.037g.al/lbSO
39^310 hrs
$ 4^938,000
j 8, 125, 000
Unit Cost
$ 0.15/KWh
$ 0.30/Mgal
_$___8/ton
te process
$ 0.015/KWh
$ 0.30/gal
•
$ ,8/hr
Annual Cost
$ 283,000
109,000
440,000
355,000
167,000
1,647,000
$ 2T169fOOO
315,000
247,000
406,000
412,000
^
i
.^
.-
^
•
^
^
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
a. Sulfuric Acid :
b. Alternative
$3,989,000
470,000
c. Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost (basis $20,400,000)
8. Total Net Annualized Cost
9. Annual Cost/ton S02 Removed
342
-------
WEAj^S02 STREAM CONTROL COPPER SMELTERS APPend*x
CAPITAL INVESTMENT COSTS
SMELTER; Kennecott - Hurley, New Mexico
CONTROL PROCESS. Citrate Process
JASIS: Maximum gas flow 270,000 SCFM
Average S02 rate 2Q.200 ihs/hr
Temperature at control point 425 °p
SPECIAL CONDITIONS
ESP cannot handle present fas flows and efficiency is poor
COST DETERMINATION
1. Gas conditioning cost « 4.40Q.QQQ
(a) Adjusted cost - 3,950,000
2. S02 Absorption cost(includes 5% - 2.450.000
TOTAL allowance) m
2
(b) H2S04 m N/A
(c) Modifications to existing »
HSO plant
6. Special costs
(a) ESP (not chargeable to process) » (4,000,000)
(b) Alternate processing equip. -
(a) Retrofit allowance ('25%C.I.) * 1.600,000
TOTAL " " 8,OOO.QQO
3. S02 Handling Section cost(19,2001b/h*0 6,200,000
(a) Disposal - 100.000
TOTAL - ' 6,300,000
4. Auxiliary plant
(a) Liquid SO,
TOTAL BOUNDARY LIMIT COSTS - 14,300,000
5. Support Services
(a) Power ( KVA) -
(b) Steam ( Ibs/hr). -
(c) Water ( 3.2Mgal/min) - 50,000
TOTAL^P^"8 °nly> - 50,000
(c) General site costs 20% item 4 - 2,860,000
TOTAL - 2.860.000
TOTAL CAPITAL INVESTMENT » 17,210,000
343 ~
-------
WEAK SO STRUM CONTROL COPPER SIMI/TICKS
ANNUAL 'OPERATING COSTS
Smelter: Kennecott - Hurley, New Mexico Control Process:
Basis:
Appendix G
Citrate
Max. Gas Flow
Av. SO- Rate_
275.00QSCFM Teinp. at Control Point 425°F
20,200 Ibs/hr
COST COMPUTATION;
1. Gas Conditioning & S0_
Absorption
a. Power [7x(|i£-)3 + 4.0 ]
1060'
b. Make-up water [3. OX(T^T)-6]
lUoU
c. Neutral. Limestone
TOTAL
2. S02 Handling (19,200 Ib/hr)
a. Chemicals
b. Power
c. Steam
d. Fuel Oil or Nat. Gas
e. Water (a) Process
(b) Cooling
f. Disposal
TOTAL
3. Labor (4 3/4 men/shift)
4. Maintenance
a. Gas Condn. @ 5 % TCI
b. S02 Handling @4 % TCI
5. Insurance & Taxes @ 2 1/2% TCI $17>210>00°
Basis
8,4/KW/MSCEM
2.7gal/MSCFM
0.091bs/MSCFW
See appropria
section
0.075KWh/lbSO
itACFj^y^Qo
l.-JM}JJ£Sp_^
li£siy.A!iso_3_
4JL,_61p___hrs
$ 4,940,000
$__9_,_2_6_0_,_0_0_0__
?17,210,000
Unit Cost
$ 0.1S TOJh
$. 0.30/Mgal
$ 8/ton
te process
o 1 0.015/KWh
i_l.._25/MCF_
$__P_._3/_Mga.L.
l_0_._l/Mgal.
$__8/h_r_
Annual Cost
$ ?fti nnn
__ —log^ooo.
48,000
440,000
563JDOO
1 176JDOO
_____J_UJIOO
7_1AOQO_
78^000.
1,601,000
______333juP_QP_
247_J)00
_3_7_°j_°_QP_
430,000 _
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
a.
b. Alternative
N/A
c. Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost
8. Total Net Annualized Cost
9. Annual Cost/ton SO- Removed
344
-------
WEAKSO, STR^j:0_NTROL__Cp_PPER SMET..TERS Appendix G
.z
CAPITAL- 1 NVF.S TMKNT COSTS
SMELTER: Kennecott-McGill, Nevada
CONTROL PROCESS Double -Alkali
BASIS:
Maximum gas
Average SO*
Temperature
flow
rate
at control
250.
20.
point
000
000 <
450
SCFM
Ibs/hr
°F
SPECIAL CONDITIONS
Existing ESP's old and recovery efficiencies low. Replacement 1? proposed.
COST DETERMINATION
1. Gas conditioning cost = 2,60Q,QQO
(a) Adjusted cost ~ 2,370,000
2. S07 Absorption cost (includes 5% = 2,360,000
allowance) _ .
TOTAL i»
6. Special costs
(a) ESP * ~
. (b) Alternate processing equip. -
(c) General site costs@20% item 4 - 3.450,000
(a) Retrofit allowance (30%C.I.) = 1.420,000
TOTAL = 6.150.000
3. S02 Handling Section cost (19,0001b/hr). 6.100.000
(a) Disposal . ~ —
TOTAL - 6,100,000
4. Auxiliary plant
(a) Liquid S02 =
(b) H2S04 = ^
(c) Modifications to existing =
H_SO, plant
TOTAL BOUNDARY LIMIT COSTS = 12,250,000
5. Support Services
(a) Power (j4000_KVA) - SBOjOpO
(b) Steam ( Ibs/hr) = —
(c) Water ( 0. 7 Mgal/min)
780.000
TOTAL ~
2 Asn nnn
TOTAL
$15,480,000
TOTAL CAPITAL INVESTMENT - .
345
-------
WEAK S02 STREAM CONTROL COPPER SMELTERS
ANNUAL "OPERATING COSTS
Appendix G
Smelter: Kennecott - McGill, Nevada
Basis: Max. Gas Flow 250,000 SCFM
Control Process: Double Alkali
Temp, at Control Point 450 F
Av. SO- Rate_
20,000 Ib/hr
COST COMPUTATION;
1.
2.
3.
4.
5.
Gas Conditioning & S0_
Absorption
a Power f5.25f ) + 4 Ol
b. Make-up water r-j/-910 x f.-i
c. Neutral. Limestone
TOTAL
S02 Handling(19,000 Ib/hr)
a. Chemicals
b. Power
c. Steam
d. Fuel Oil or Nat. Gas
e. Water
f. Disposal
TOTAL
Labor (3 men/shift)
Maintenance
a Gas Condn. @ ^ % TCI
b SO Handling @ 4 % TCI
Insurance & Taxes @ 2 1/2% TCI
Basis
7.3KW /MSCFM
2.7gal/MSCFMj
0.071b /MSCFM j
See appropria
section
0.075KWh/lbSO
a^gaLOhSO^..
ajLUh/JfcSQg..
26i28p_jLrs.__.
lALflfiJUflfifl. .
?15, 480, 000
Unit Cost
$ 0.015/KWh
$ 0.3/Mgal
$ 8/ton
te process
2_ $0.015/KWh
-ULJl/JIgai.
Annual Cost
223,000
99,000
34,000
356.000
2,400^000
174,000
Q nnn
821 QOA
3,404,000
_2ifumo_
J6J_QflQ_
387,000 __,,
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
4,878,000
a.
b. Alternative
c. Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost
8. Total Net Annualized Cost
N/A
$4,878,000
$2,036,000
$4,n77,nno
9. Annual Cost/ton SO. Removed
346
$53
-------
Appendix G
>ER SMELTERS
CAPITAL INVESTMENT COSTS
WEAK S02 STREAM CONTROL COPPER SMELTERS
SMELTER;
CONTROL PROCESS. Magnesium Oxide
BASIS; Maximum gas flow 250,000 SCFM
Average SO- rate :.20;i,000 ibs/hr
Temperature at control point 450 p
SPECIAL CONDITIONS
Existing ESP's old and recovery efficiencies low. Replacement is proposed.
COST DETERMINATION
4,150,000
1. Gas conditioning cost
(a) Adjusted cost = 3,790,000 _
2. S02 Absorption cost (includes 5% - 2,680,000
TOTAL allowance) = 6.470.000
(a) Retrofit allowance (30% C.I.) = 1.940.000
- TOTAL - 8.410.000
3. S02 Handling Section cost - 4.600.000
(a) Disposal (18,000 Ib/hr) - HII
TOTAL - 4.600.000
4. Auxiliary plant
(a) Liquid S02 " —
(b) H2S04 ' —=
(c) Modifications to existing -
H2SO, plant
TOTAL BOUNDARY LIMIT COSTS - 16.960.000
5. Support Services
(a) Power (6500 KVA) - ^0,000
(b) Steam ( Ibs/hr).
6. Special costs ~~
. (a) ESP "
(b) Alternate processing equip. " „
(c) General site costs « 20% item 4 . 3.390.000
(c) Water ( 1.2 Mgal/min) - 550^000
i.ooo.ooo
TOTAL "
3,390,000
TOTAL "
$21,350,000
TOTAL CAPITAL INVESTMENT " —
347
-------
WEAK S02 STREAM CONTROL COPPER SMELTERS
ANNUAL OPERATING COSTS
Smelter: Kennecott-McGill, Nevada Control Process:
rt f f\ f\f\f\e* s*-mr
Basis: Max. Gas Flow
Appendix G
Oxide
250.000SCFM
Temp, at Control Point
450°F
Av. SO- Rate_
20,000 Ib/hr
COST COMPUTATION;
1. Gas Conditioning & SO.
Absorption
-6]
a. Power L
b. Make-up water[
c. Neutral. Limestone
TOTAL
2. S02 Handling (18,000 Ib/hr)
a. Chemicals
b. Power
c. Steam
d. Fuel Oil or Nat. Gas
e. Water
f. Disposal
TOTAL
3. Labor ^< man/shif$ + 1 day)
4. Maintenance
a. Gas Condn. @ % TCI
b. S02 Handling
-------
Appendix G
WEAK S02 STREAM CONTROL COPPER SMELTERS
CAPITAL INVESTMENT COSTS
SMELTER: Kennecott-McGill, Nevada
CONTROL PROCESS. Citrate
BASIS; Maximum gas flow 250,000 SCFM
Average S02 rate 2,0.. OOP- ibs/hr
Temperature at control point Asn F
SPECIAL CONDITIONS
Existing ESP's old and recovery efficiencies low. Replacement is proposed.
COST DETERMINATION
1. Gas conditioning cost = 4.150.000
(a) Adjusted cost - 3.790.000
2. SO, Absorption cost(includes 5% - 2,360,000
2 allowance) , , cn nnn
TOTAL * 6,150,000
(a) Retrofit allowance (30% C.I.) = 1.840,000
TOTAL - 7."0'000
3. SO Handling Section cost(19,000 Ib/hc) 6,200,000
(a) Disposal - 100>°00
6. Special costs
. (a) ESP
(b) Alternate processing equip.
TOTAL - 6.300,000
4. Auxiliary plant
(a) Liquid S02
(b) H2S04
(c) Modifications to existing -
H2SO, plant
TOTAL BOUNDARY LIMIT COSTS - 14.290,000
5. Support Services
(a) Power (420C
(b) Steam ( Ibs/hr). -
(c) Water (_l^_Mgal/tain) - 600»000
TOTAL - 98Q.QQQ
(c) General site costs @20% item 4 - 2.820.000
TOTAL " 2.820.000
TOTAL CAPITAL INVESTMENT - $18.090.000
349
-------
WEAK S02 STREAM CONTROL COPPER SMELTERS Appendix G
ANNUAL OPERATING COSTS
Smelter; Kennecott - McGill. Nevada Control Process! Citrate
250,000 SCFM
Basis:
Max. Gas Flow
Av. SO- Rate
Temp, at Control Point
20,000 Ib/hr
COST COMPUTATION;
1. Gas Conditioning & S02
Absorption
a. Power
b. Make-up water
c. Neutral. Limestone
TOTAL
2. S02 Handling (19,000 Ib/hr)
a. Chemicals
b. Power
c. Steam
d. Fuel Oil or Nat. Gas
e. Water (a) Process
f. Disposal Cooling
TOTAL
3. Labor (4 3/4 man/slhift) 3
4. Maintenance
a. Gas Condn. (?__5_Z TCI
b. S0_ Handling
-------
WEAK S02 STREAM CONTROL COPPER SMELTERS : Appendix G
CAPITAL INVESTMENT COSTS
SMELTER; Magma Copper - San Manuel, Arizona
CONTROL PROCESS Double Alkali
BASIS; Maximum gas flow '_ SCFM
Average S02 rate :34,:500 Ibs/hr
Temperature at control point 500 F
SPECIAL CONDITIONS
COST DETERMINATION
1. Gas conditioning cost = 4.200,000
(a) Adjusted cost - 3.950,000
2. S09 Absorption cost(includes 5% = 3.470.000
TOTAL allowance) = 7.420,000
(a) Retrofit allowance (25%C.I.) = 1.860,000
TOTAL " 9.280.000
3. S02 Handling Section cost(32,8001b/h*) 8,700.000
(a) Disposal
TOTAL - 8'700-000
6. Special costs
. (a) ESP "
(b) Alternate processing equip. " _—.
(c) General site costs 15% item 4 . 2,700,000
4. Auxiliary plant 9
(a) Liquid S02 = -
(b) H2S04 " - — -
(c) Modifications to existing - -
H2SO^ plant
TOTAL BOUNDARY LIMIT COSTS - 17.980,000
5. Support Services
(a) Power (9000 KVA) - _ 520,000 -
(b) Steam ( _ Ibs/hr)
(c) Water (J^7_Mgal/min) - 640,000
1,160,000
TOTAL *
2,700,000
TOTAL .
$21,840,000
TOTAL CAPITAL INVESTMENT -
351
-------
WEAK S02 STREAM CONTROL COPPER SMELTERS Appendix G
ANNUAL 'OPERATING COSTS
Smelter: Magma Copper - San Manuel. Arizona Control Process.' Double Alkali
Basis:
Max. Gas Flow
550,000 SCFM
Temp, at Control Point 500 F
Av. SO- Rate_
34.500 Ib/hr
COST COMPUTATION:
1. Gas Conditioning & SO,
Absorption
960 ,3
a. Power[5.25(7^7T)J + 4.0]
1060'
b. Make-up water [3.0(|^> '6]
c. Neutral. Limestone
TOTAL
2. S02 Handling (32,800 Ib/hr)
a. Chemicals
b. Power
c. Steam
d. Fuel Oil or Nat. Gas
e. Water
f. Disposal
TOTAL
3. Labor (3 man/shift) 3
4. Maintenance
a. Gas Condn. @ 5 % TCI
b. S02 Handling @ 4 % TCI
5. Insurance & Taxes @ 2 1/2% TCI
Basis
7 . 9KWh/MSCFM
2.8gal/MSCFM
0.071b/MSCFM
See appropria
section
0.075KWh/lbSC
0.2gal/lbSO
3.531b/lbS00
26,280 hrs
$ 4,938,000
$13,038,000
$21,840,000
Unit Cost
$ 0.015/KWh
$ 0.3/Mg^al
$ 8/ton
te process
~ $ 0.015/KWh
$ 0.3/Mgal
$ 3/ton
$ 8/hr
Annual Cost
532JJOO
226,000
75,000
833,000
4,142,000
301,000
16,000
1,417,000
5,876,000
210,000
247,000
522,000
546,000 ^^
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
8.234.000
a.
b. Alternative
N/A
c. Incremental Costs Associated with Use
of. Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost
8. Total Net Annualized Cost
9. Annual Cost/ton SO- Removed
?8.234.OOQ_
$2,872,000.
$6.455,000.
$48
352
-------
WEAK S02 STREAM CONTROL COPPER SMELTERS
CAPITAL INVESTMENT COSTS
SMELTER; Magma Copper - San Manuel, Arizona
CONTROL PROCESS. Magnesium Oxide
BASIS; Maximum gas flow • 550,000 gCFM
Average S02 rate 34»:5.9°.. Ibs/hr
Temperature at control point 500 F
SPECIAL CONDITIONS
Appendix G
COST DETERMINATION
1. Gas conditioning cost - 7.000.000
(a) Adjusted cost - 6,580.000
2. SO, Absorption cosCUncludes 5% - 3»990>000
TOTAL allowance) » 10.570,000
(a) Retrofit allowance (25% C.I.) = 2.640,000
TOTAL * " 13.210.000
3. S02 Handling Section cost (31,0001b/h*) 6.600.000
(a) Disposal =
TOTAL - 6.600.000
5. Support Services
(a) Power £0.20QKVA) - 550^000.
(b) Steam ( Ibs/hr). - _
TOTAL
6. Special costs
' (a) ESP
(b) Alternate processing.equip.
(c) General site costs @15Z item A - 3,820,000
TOTAL
Auxiliary plant
(a) Liquid SO, *
, * - ' 5,650,000
(b) H2S04 - '
(c) Modifications to existing -
H2SO^ plant
„„,.... 25,A60,000
TOTAL BOUNDARY LIMIT COSTS » -—.—.2...—..-—-
(c) Water (1.8 Mgal/min) = 680,000
1.230.000
TOTAL CAPITAL INVESTMENT - $30,510.000_
353
-------
WEAK SO. STREAM CONTROL COPPER SMELTERS
Appendix G
ANNUAL OPERATING COSTS
Smelter: Magma Copper - San Manuel, Arizona ' Control Process: Magnesium Oxide
Basis:
Max. Gas Flow
Av. SO. Rate_
550,000 SCFM
Temp, at Control Point_
500°F
34,500 Ib/hr
COST COMPUTATION:
1. Gas Conditioning & S02
Absorption
+ 4.0]
.960 , t
a. Power L'.(
b. Make-up water [3.0<
c. Neutral. Limestone
TOTAL
2. S02 Handling (31,000 ib/hr)
a. Chemicals
b. Power
c. Steam
d. Fuel Oil or Nat. Gas
e. Water
£. Disposal
TOTAL
3. Labor (4*4 men/shift + 1 day) 3
4. Maintenance
a. Gas Condn.
-------
Appendix G
WEAK SO,. STREAM CONTROL COPPER SMELTERS
CAPITAL INVESTMENT COSTS
SMELTER: Ma&ma Copper - San Manuel, Arizona
CONTROL PROCESS. Citrate
BASIS; Maximum gas flow _ 550,000 _ SCFM
Average S02 rate _ ^.ftPO^ Ibs/hr
Temperature at control point _ 5(30 _ Op
SPECIAL CONDITIONS
COST DETERMINATION
1. Gas conditioning cost - 7,000,000
(a) Adjusted cost - 6,580,000
2. S0_ Absorption cost( includes 5% - 3,470,000
allowance) in ncn nnn
TOTAL * 10,050,000
(a) Retrofit allowance(25% C.I.) = 2,510.000
TOTAL - 12.560.000
3. S02 Handling Section cost (32,8001b/h») 8.300,000
(a) Disposal - 200»000
6. Special costs
(a) ESP
(b) Alternate processing equip,
TOTAL - 8.500.OOP
4. Auxiliary plant
(a) Liquid S02 "
(b) H2SOA - ^
(c) Modifications to existing »
H_SO, plant
TOTAL BOUNDARY LIMIT COSTS - $21.060.000
5. Support Services
(a) Power (9500 KVA) - 53°'°00
(b) Steam ( Ibs/hr). - -
(c) Water ( 2.6 Mgal/min) - 80M°0
1,330,000
TOTAL "
(c) General site costs (§15% item 4 - 3.160.000
TOTAL - 3,160,000
TOTAL CAPITAL INVESTMENT - ' $25,550,0j)p_
355
-------
Appendix G
WEAK S02 STREAM CONTROL COPPER SMELTERS
ANNUAL OPERATING COSTS
Smelter; Magrca c°PPer ~ San
Basis: Max. Gas Flow
' Arizona
control Process: Citrate
550.000SCFM
Temp, at Control Point_
Av. S02 Rate_
34,500 Ib/hr
COST COMPUTATION;
1. Gas Conditioning & S02
Absorption
a. Power
b. Make-up water
c. Neutral. Limestone
TOTAL
2. S02 Handling (32,800 ib/hr)
a. Chemicals
b. Power
c. Steam
d. Fuel Oil or Nat. Gas
e. Water (a
^ Process
(b) Cooling
TOTAL
3. Labor (4 3/4 man/Shift) J
4. Maintenance
a. Gas Condn. @_5_Z TCI
b. SO- Handling @ 4 Z TCI
£. "
5. Insurance & Taxes @ 2 1/2Z TCI
Basis
9.2KWh/MSCFM
2.8gal/MSCFM
0.091b/MSCFM
See appropria
section
'8,225,000
25,550,000
Unit Cost
0.015/KWh
i_ _°_-3/MJal
$ 8.0/ton
e process
_^_8/hr
Annual Cost
6_1_9_,000_
J.26_,000_
97,000
942,000
963,000
_2Q_1_Q-OJL
1^204^000 '
134^000
2,722,000
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
a.
b. Alternative
c. Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost
8. Total Net Annualized Cost
9. Annual Cost/ton SO- Removed
356
-------
WEAK SO STREAM CONTROL COPPER SMELTERS
ANNUAL 'OPERATING COSTS
Phelps Dodge-Ajo, Arizona
Smelter:
Basis: Max. Gas Flow
55,000 SCFM
Appendix G
. *Dimethylaniline
Temp, at Control Point 450°F
Av. S0_ Rate_
10,100 Ibs/hr
*Note: Includes S0» liquifaction
COST COMPUTATION;
1. Gas Conditioning
,910
a. Power
r7,
>- (
b. Make-up water
c. Neutral. Limestone
TOTAL
2. SO- Absorption & Handling
L (9,700 Ib/hr or 2.0%)
a. Chemicals
(1) Absorption
b. Power
(2) S02 Handling
c. Stream
d. Water
TOTAL
3. Labor (2 man/shift)
4. Maintenance
a. Gas Condn. @_5 _% TCI
b. S02 Handling @£_% TCI
5. Insurance & Taxes @ 2 1/2% TCI
Basis
4.4KW/MSCFM
2.7gal/MSCFM
0.091b/MSCFM
See appropria
section
8.0KW/MSCFM
0.02KWh/lbSO.
1.51b/lbSO%
42.4gal/MSCF*
17,520 hrs
$ 1,630,000
$ 7,200,000
$ 8,830,000
Unit Cost
$ 0.015/KWh
$ 0.30/Mgal
$ 8/ton
te process
$_ 0.015/KWh
$ 1.25/Mlbs
$ 0.10/Mgal
$ 8/hr
Annual Cost
30,000
22,000
10,000
62,000
4_92_J1Q_0_
54_J300
_— ?AiOOO
148J)00
114,000
832,000
140,000
81,000
288_,000
220,000
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
1.623.000
b. Alternative
N/A
c. Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7- Annualized Capital Coat
8- Total Net Annualized Cost
9< Annual Cost/ton SO, Removed 357
$1.623.000
$1.161,000
$1,722,000
$43
-------
WEAK SO. STREAM CONTROL COPPER SMKf.TKRS
. . . ^ .._. ._- ._
CAIMTAI-, INVI'STMKNT COSTS
SMELTER: Phelps-Dodge - Douglas, Arizona
CONTROL PROCESS Citrate Process
Roaster/Reverb Converter
BASIS: Maximum gas flow 445.000 SCFM 265,000 SCFM
70,100 Ibs/hr
o,.
Average S02 rate
43.000 Ibs/hr
Temperature at control point
SPECIAL CONDITIONS
410
350 °F
Two separate gas conditioning and SO,, absorption systems with a common SO,
"•• - - "" ~ '~ ' _ - -I - - - - ^T • Tr—.1 .----- - ^
regeneration section
6,100,000
5,430,000
3,000,000
COST DETERMINATION
1. Gas conditioning cost
(a) Adjusted cost
2. S0? Absorption cost
TOTAL
(a) Retrofit allowance <50%c.l.)
TOTAL
3. S02 Handling Section cost(107,4001b/far)16,000,000
8,430,000
4,220,000
250,000
(a) Disposal
TOTAL
4. Auxiliary plant
(a) Liquid S02
(b) H2S04
(c) Modifications to existing =
H SO, plant
TOTAL BOUNDARY LIMIT COSTS =
5. Support Services
(a) Power IJJ^OOpKVA)
(b) Steam ( Ibs/hr)
(c) Water ( 5.0 Mgal/min) =
TOTAL
6. Special costs
(a) ESP =
(b) Alternate processing equip. =
(c) General site costs (15% item 4)= 5,870,000
TOTAL
TOTAL CAPITAL INVESTMENT
650,000
1,250,000
4.350,000
3,700,000
2,330.000
6,030,000
(70%C.I.) 4,220,000
$22,900,000
16.250.OQq
39.150.000
1.900.000,
5.870.000.
358
-------
WEAK SO. STREAM CONTROL COPPER SMELTERS
ANNUAL OPERATIUG COSTS FOR GAS CONDITIONING AMD S02 ABSORPTION
Smelter: Phelps-Dodge-Douglas, Arizona
Control Process: Citrate
A. Roaster/Reverb Gas Stream
Basis: Max. Gas Flow 445.000 SCFM
Av. S0? Rate_
70,100 Ib/hr
COST COMPUTATION;
1. Gas Conditioning &
Absorption
a. Power
+ 4]
b. Make-up water 870 \.
c. Neutral. Limestone
TOTAL
Temp, at Control Point
410°F
Basis
7.9KW/M_SC_FM_
Unit Cost
_$0.0157KWh
^O.JI/Mgal
Annual Cost
.^aOj.ooo,.
-Iie^QQQ..
.-28^000—
684,000
B. Converter Gas Stream
Basis:
Max. Gas Flow
Av. SO- Rate_
COST COMPUTATION;
1. Gas Conditioning & S02
Absorption
a. Power
b. Make-up water -,,810
1 Of
c. Neutral. Limestone
TOTAL
SCFM
Temp, at Control Point
43.000 Ib/hr
Basis
7.1KW/MSCFM
2.6gal/MSCFM
D.091b/MSCFM
Unit Cost
SP.OlSj/KWh
|0.3^Mgal
J8j/ton ._.
Annual Cost
_ _ 23Q..QQO-
1QUQQQ _
4Z^QQQ
378,000
GRAND TOTAL
$1.062.000
359
-------
WEAK SO- STREAM CONTROL COPPER SMI1LTERS
ANNUAL 'OPERATING COSTS
Smelter: Phelps-Dodge - Douglas, Arizona Control Process: Citrate
Basis: Separate Gas Handling Facilities - Central S0? Regeneration
_. 113.100 Ib/hr
COST COMPUTATION:
Total Av. SO- Rate_
1. Gas Conditioning & SO,
Absorption
Total for Both
Gas Systems
TOTAL
2. S02 Handling (107,400 Ib/hr)
a. Chemicals
b. Power
c. Steam
d. Fuel Oil or Nat. Gas
e. Water (a) Process
(b) Cooling
TOTAL
3. Labor 8 man/shift
4. Maintenance
a. Gas Condn. @ 5 % TCI
b. S02 Handling @ 4% TCI
5. Insurance & Taxes @ 2 1/2% TCI |$46 ,920_,JiQJ
Basis
See appropri
section
0..075KWh£lbSC
3.6CF/lbSO,
i:JL&al£lbSO.,
70_2_080_hr_s
$Iiu435_up_00
$24,461,000
Unit Cost
te process
_$_8/_hr.
Annual Cost
1,062,000
3,152,000
8,914,000
722.000
i.m,onn
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
a.
13,410,000
b. Alternative
c
MA.
Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost
8. Total Net Annualized Cost
9. Annual Cost/ton S0_ Removed
$13,410,000_
$ 6,170,000
$11.642.000-
$27
360
-------
WEAK S_Q2 STREAM CONTROL COPPER SMELTERS • Appendix G
CAPITAL- INVESTMENT COSTS
SMELTER: Phelps Dodge - Morenci, Arizona
CONTROL PROCESS Double Alkali
BASIS; Maximum gas flow 500,000 SCFM
Average S02 rate 43,000 Ibs/hr
Temperature at control point 450 F
SPECIAL CONDITIONS
Extremely congested operating area - limited space availability
COST DETERMINATION
1. Gas conditioning cost = 4,100.000
(a) Adjusted cost = 3.690.000
3,300,000
2. S09 Absorption cost (includes 5/t =
1 allowance)
TOTAL = 6,990,000
(a) Retrofit allowance (70%C.I.) = 4,890.000
TOTAL = 11.880.000
3. SO. Handling Section cost = 10.000,000
2 (40,000 Ib/hr)
(a) Disposal
TOTAL =
4. Auxiliary plant
(a) Liquid S02 "
(b) H2S04 = .
(c) Modifications to existing =
H0SO, plant
24^
TOTAL BOUNDARY LIMIT COSTS = 21,880,000
5. Support Services
(a) Power (8000 KVA) - 520»000
(b) Steam ( Ibs/hr)
(c) Water (_1_5 M^al/mln) - IQPjJLQP.
TOTAL " " - 1,120.000
6. Special costs
. (a) ESP =
(b) Alternate processing equip. ** _
(c) General site costs!? 20% Item 4 = 4,370,000
TOTAL " . 4,370,000
$27,370,000
TOTAL CAPITAL INVESTMENT - . ===========
361
-------
WEAK K02 STREAM CONTROL COPPER SMELTERS
Appendix G
ANNUAL 'OPERATING COSTS
Smelter: Phelps Dodge - Morenci, Arizona Control Process; Double Alkali
Basis: Max. Gas Flow 500,000 SCFM
Temp, at Control Point 450°F
Av. SO- Rate
43,000 Ib/hr
COST COMPUTATION;
1. Gas Conditioning & S0?
Absorption
3
a. Power
1060
b. Make-up water
r3f 910.
[•HTnZ/vJ
1060
c. Neutral. Limestone
TOTAL
2. S02 Handling (40,800 Ib/hr)
a. Chemicals
b. Power
c. Steam
d. Fuel Oil or Nat. Gas
e. Water
f. Disposal
TOTAL
3. Labor 3 men/shift *
4. Maintenance
a. Gas Condn. @ 5_" 7CI
b. S02 Handling 04" TCI
5. Insurance & Taxes '• 2 1/2% TCI
Basis
7.3 KW/mSCFM
2.7_gal/mSCFM
0.07 Ib/mSCFM
See appropria
0.075kwh/lbSO
0.2 g_al/lbS02
3.53 Ib/lbSO^
26,280 hrs
$ 6,273,000
$15,6-10,000
?27,370.000
Unit Cost
$_0.0157kwh_
$0.3/mgal
$8/ton
2_ $0.015/kwh
-$P_, 3/jngal.
J37j:uiu
$8/hr
Annual Cost
$ 447^000
198,000
68,000
713,000
5,003,000
375,000
20,000
1,763,000
7,161,000 ___=*=
210,000
314,000
624,000
TOTAL ANNUAL OPERATING CCST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
9,706,000
a.
b. Alternative
N/A
c. Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost
8. Total Net Annualized Cost
9. Annual Cost/ton SO- Removed
362
-------
J>P_2 STREAM CONTROL COPPER SMELTERS
CAPITAL- INVESTMENT COSTS
SMELTER: Phelps Dodge - Morenci, Arizona
CONTROL PROCESS Citrate
BASIS; Maximum gas flow 500.000 SCFM
Average S02 rate 43,000 Ibs/hr
Temperature at control point 450 F
SPECIAL CONDITIONS
Extremely congested operating area - limited space availability
Appendix G
COST DETERMINATION
1. Gas conditioning cost = 6,500,000
(a) Adjusted cost - 5,850,000
2. S02 Absorption cost (includes 5% = 3,300,000
TOTAL allowance) = 9,150,000
(a) Retrofit allowance (70ZC.I.) - 6,400.000
TOTAL = 15,550,000
3. S09 Handling Section cost (40,800 lb/hrXJi:*°°»°2°
, r = 200,000
(a) Disposal — —
TOTAL " 9,600,000
4. Auxiliary plant
(a) Liquid S02
(b) H2S04
(c) Modifications to existing =
H0SO, plant
2 H
TOTAL BOUNDARY LIMIT COSTS - 25,150,000
5. Support Services
(a) Power (IQ.OOfrVA) - 56^000
(b) Steam ( Ibs/hr)
(c) Water (_2^.Mgal/min) - __800^00 ------
1.360.000
TOTAL -
6. Special costs
• (a) ESP * — - ---
(b) Alternate processing equip. " - _ - -
(c) General site costs @ 20% Item 4= 5>000'00° --
5.0QO.QQQ
TOTAL —
TOTAL CAPITAL INVESTMENT - $31.510,000_
363
-------
WRAK KC>2 STREAM CONTROL COPPER SMELTERS
Appendix G
ANNUAL 'OPERATING COSTS
Smelter: Phelps Dodge - Morenci, Arizona Control Process; Citrate
Basis:
Max. Gas Flow
Av. SO Rate_
500,000 SCFM
Temp, at Control Point 450°F
43,000 Ib/hr
COST COMPUTATION;
1. Gas Conditioning & SO-
Absorption
a. Power [7^) +4]
b. Make-up water U/ 910v " I
LJ4060; -I
c. Neutral. Limestone
TOTAL
2. S02 Handling (40,000 Ib/hr)
a. Chemicals
b. Power
c. Steam
d. Fuel Oil or :;=t. Gas
e. Water (a) Process
XXXXEXXJQSJSX2 (b) Cooling
TOTAL
3. Labor (4.75 man/shift)
4. Maintenance
a. Gas Condn. I: 5 % TCI
b. S02 Handling 'i 4% TCI
5. Insurance & Ta~e = 3 2 1/2% TCI
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Planr
a.
b. Alternative
c. Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost
8. Total Net Annualized Cost
9. Annual Cost/ton SO- Removed
Basis
8.4 kw/mSCFM
2.7 gal/mSCFM
0.09 Ib/mSCFM
5ee ajpjprojpria
3.075kwh/lbSO
3.6 CF/lbS02_
L.5 gal/lbS02
5.0 gal/lbS02
41,610 hrs
$ 9,945,000
$15,210,000
$31,510,000
Unit Cost
$0.015/kwh
$0.3/mgal
$8/ton
:e process seci^
|_$0.015/kwh
$1.25/MCF
$0.3/mgal
$0.1/mgal
$8/hr
Annual Cost
$ 514,000
198,000
88,000
800,000
1,198,000
375,000
It498j)00
ISOJIOO
166^000
3,387,000
333,000
497,000
608,000
788,000 _—
6.413,000
N/A
6,413,000
4,144,000
6,470,000
364
$39
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WEAK S02 STREAM CONTROL COPPER SMELTERS • Appendix G
CAPITAL INVESTMENT COSTS
SMELTER: White Pine - White Pine, Michigan
CONTROL PROCESS Double Alkali
BASIS; Maximum gas flow 110.000 SCFM
Average S02 rate 18.500' :> Ibs/hr
Temperature at control point 550 °F
SPECIAL CONDITIONS
This is a converter gas stream which operates only approximately 9-10 hours/dav.
SO;? rate has been averaged over 24 hours.
COST DETERMINATION
1. Gas conditioning cost = 1.670,000
(a) Adjusted cost = 1.620,000
2. S02 Absorption cost (includes 7% = 1.530,000
TOTAL allowance) = 3,150,000
(a) Retrofit allowance (15%C.I.) = 470,000
TOTAL = 3.620.nnn
SO, Handling Section cost = 5,800,000
, t (17,500 Ib/hr)
(a) Disposal "
TOTAL - 5.800.000
4. Auxiliary plant
(a) Liquid S02 -
Support Services
(a) Power (2600 KVA) - 320,000
(b) Steam ( Ibs/hr). -
(c) Water ( Mgal/min) *
6. Special costs_
• (a) ESP
(b) Alternate processing equip.
(b) H2S04 = N/A
(c) Modifications to existing =
H2SO/, plant
TOTAL BOUNDARY LIMIT COSTS • 9.420.000
TOTAL * 320,000
(c) General site costs (20% item 4)- 1,880.000
TOTAL " 1.880.000
TOTAL CAPITAL INVESTMENT - 11.620,000^
365
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WEAK S02 STREAM CONTROt COPPER SMELTERS Appendix G
ANNUAL OPERATING COSTS
Smelter; White Pine - White Pine, Michigan * Control Process; Double Alkali
550°F
Basis:
Max. Gas Flow 110.000 SCFM
Av. S0_ Rate 18,500 Ibs/hr
Temp, at Control Point_
COST COMPUTATION;
1. Gas Conditioning & S0_
Absorption ^ o
a. Power
b. Make-up water Ig/
+ 4.0~|
Basis
Unit Cost
060
Neutral. Limestone
TOTAL
2. S02 Handling (17,500 Ib/hr)
a. Chemicals
b. Power
c. Steam
d. Fuel Oil or Nat. Gas
e. Water
f. Disposal
TOTAL
3. Labor (3 men/shift)* * *
4. Maintenance
a. Gas Condn. @ 5 % TCI
b. SO- Handling @ 4 Z TCI
^
5. Insurance & Taxes Q 2 1/2Z TCI
0.07 Ib/mSCFM $8/ton
cop_ria ;e process_ sect_.__
0.075 kwh/lb 902 $0.3/iagal
™""~^^™™~™"™"^1 ' ""^••••i •"•**! i «^ i i ii ••^••^••••^•
[h_2_gal/lb Sp2|_$p_.3/'ia gal
3.53 Ib/lb SOo $3/ton . .
26,280 hrs
1$ i._8_6_3_,000_
$ 7,560,000
$11,620,000
_$8/hr_
Annual Cost
_$___ii4ioqo
47,000__
15,000
176,000
161,000
_ 9,000
756,000
3,137,000
210,0_qp_
_93,000
302,000
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
a.
b. Alternative
c. Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost
8. Total Net Annualized Cost
N/A
$ 4.209.000,
$ 1,528,000
$ 3,345,000,
9. Annual Cost/ton SO- Removed
366
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Appendix G
WEAK S02 STREAM CONTROL COPPER SMELTERS
CAPITAL INVESTMENT COSTS
SMELTER: White Pine - White Pine, Michigan
CONTROL PROCESS Magnesium Oxide
BASIS; Maximum gas flow 110.000 SCFM
Average S02 rate 18,500 ',: :, lbs/hr
Temperature at control point 550 p
SPECIAL CONDITIONS
This is a converter gas stream which operates only approximately 9-10 hours/day.
S02 rate has been averaged over 24 hours.
COST DETERMINATION
1. Gas conditioning cost - 2.450,000
(a) Adjusted cost - 2,380,000
2. SO. Absorption cost (includes 7% - 1,720,000
TOTAL allowance) = 4.100.000
(a) Retrofit allowance (15%C.I.) - 620.000
TOTAL * 4.720.000
3. S00 Handling Section cost (16,600 . - 4,400.000
, J lb/hr)
(a) Disposal
TOTAL * ', 4.400.000
TOTAL
4 . Auxiliary plant
(a) Liquid S02 • -
(b) H2S04 - - 3,750,000
(c) Modifications to existing = -
H2SO, plant
TOTAL BOUNDARY LIMIT COSTS - 12,870,000
5. Support Services
(a) Power ( 3200 KVA) - 35Q»000 - -
(b) Steam ( _ lbs/hr). - - _ -
(c) Water ( _ Mgal/rain) • - : -
TOTAL «
6. Special costs
• (a) ESP - " - - - -
(b) Alternate processing equip. a __ - .
(c) General site costs (20% item 4)- 2.570,000 -
2,570.000
TOTAL CAPITAL INVESTMENT - ======_==
367
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WEAK S02 STREAM CONTROt. COFFER SMELTERS
Appendix G
ANNUAL OPERATING COSTS
Smelter; White Pine " White Pine, Michigan ' Control Process; Magnesium oxide
550°F
Basis:
Max. Gas Flow
Av. S0_ Rate_
110.000 SCFM
18,500 Ib/hr
Temp, at Control Point
COST COMPUTATION;
1. Gas Conditioning & SO.
Absorption
b. Make-up water
c. Neutral. Limestone
TOTAL
2. S02 Handling (16,600 Ib/hr)
a. Chemicals
b . Power
c. Steam -
d. Fuel Oil or Nat. Gas
e. Water
f. Disposal
TOTAL
3. Labor (&u men/shift + 1 day) * j
4. Maintenance
a. Gas Condn. 0 5 Z TCI
b. SO- Handling 05 Z TCI
£.
5. Insurance & Taxes @ 2 1/2Z TCI
Basis
9.95 kwh/mSCI
Unit Cost
$0.015/kwh
2.9 gal/mSCFM] $0.3/mgal
0.09 Ib/mSCFM
$8/ton
See appropriate process sectio
0.1 kwh/lb SOof 40.015/kwh
0^037 gal/lb 902 $0.3/gal
0.2 gal/lb Sol $0.3/mgal
39,310 hrs
$6,378,000
$12,040,000
$8/hr
Annual Cost
$ 134,000
19,000_
200,000
324±000
203,000
1,504,000
8,000
2,039,000
315,000
137_,_000
319,000
301,000
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
a. H2S04 plant .
b. Alternative
3,311,000
450,000
c. Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost (basis $15,790,000)
8. Total Net Annualized Cost
9. Annual Cost/ton S02 Removed 368
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Appendix G
WEAK S02 STREAM CONTROL COPPER SMELTERS
CAPITAL INVESTMENT COSTS
SMELTER; White Pine - White Pine. Michigan
CONTROL PROCESS Citrate
BASIS; Maximum gas flow 110,000 SCFM
Average S0? rate 18,500 :..! :. lbs/hr
Temperature at control point 550 °p-
SPECIAL CONDITIONS
This is a converter gas stream which operates only approximately 9-10 hours/day.
S02 rate has been averaged over 24 hours.
COST DETERMINATION
1. Gas conditioning cost = 2.450.000
(a) Adjusted cost = 2.380.000
2. S02 Absorption cost(includes 7% - 1.530.000
TOTAL allowance) = 3.910.000
(a) Retrofit allowance (15ZC.I.) - 590.000
TOTAL -- 4.500.000
3. S02 Handling Section cost(17,500 = 5.900.000
(a) Disposal lb/hr) - 100.000
TOTAL * 6.000.000
4. Auxiliary plant
(a) Liquid S02 -
(b) H0SO, -
6_. Special costs
• (a) ESP -
(b) Alternate processing equip.
(c) General site costs (20% item 4). 2.100,000
(c) Modifications to existing =
H2SO, plant
TOTAL BOUNDARY LIMIT COSTS - 10,500,000
5. Support Services
(a) Power ( 3000 KVA) - 330,000
(b) Steam ( lbs/hr) - .—
(c) Water ( Mgal/min) -
TOTAL "
TOTAL - 2'100'00°
TOTAL CAPITAL INVESTMENT - 12 t93Q.QQfl.
369
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WEAK S02 STREAM CONTROt COPPER SMELTERS
ANNUAL OPERATING COSTS
Appendix G
Smelter; White Pine -'White Pine, Michigan
Basis: Max. Gas Flow
Control Process: Citrate
nh.nnn
Temp, at Control Point 550°F
Av. SO- Rate_
18,500 Ib/hr
COST COMPUTATION:
1. Gas Conditioning & SO,
Absorption
Basis
Unit Cost
2
, 7/±"i»\-'L/l • 9.95 kwh/mSCFlf $0.015/kwh $ 134,000
a. Power |>7(T7^)+M 4
J 2.9 gal/mSCFM $0.3/mgal 47,000
b. Make-up water nnn
' 0.09 Ib/mSCFM $8/ton 19,000
c. Neutral. Limestone '•
TOTAL ' 200,000
2. S02 Handling (17,500 Ib/hr)
a. Chemicals See_ap^ropjrJ._a]:e_£roces_s_jjctio _51^0QO
b. Power
_ , _,, „ _ . 3.6"cF7lbS02 $1.25/MCF 643,000
d. Fuel Oil or Nat. Gas *
. „ 1.5 gal/lb SC? $0.'3/mgal 64,000
e. Water a) Process 4'
f. BlXpSiStt b) Cooling _1Lw_^mri_'l'K. T "°~ 71,000
TOTAL I I 1^,453,000:
_ _ , *,. TC . ,_ , a, 41,610 hr $8/hr \ 333,000
3. Labor (4.75 men/shi*t) *
4. Maintenance
$ 2,737,000 137,000
a. Gas Condn. @ 5 Z TCI
$ 7,760,000 310,000
b. S02 Handling @ 4 Z TCI
!$12,930,000 323,000
5. Insurance & Taxes @ 2 1/2Z TCI
Annual Cost
TOTAL ANNUAL OPERATING COST FOR PRIMARY CONTROL PROCESS
6. Auxiliary Plant
a.
2,756,000
b. Alternative
c.
N/A
Incremental Costs Associated with Use
of Existing Acid Plant
TOTAL ANNUAL OPERATING COST
7. Annualized Capital Cost
8. Total Net Annualized Cost
2,756,000
1,700,000
2,720,000
9. Annual Cost/ton S02 Removed
370
$38
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TECHNICAL REPORT DATA
(Please read Instructions on the reverse before completing)
3. RECIPIENT'S ACCESSION-NO.
1. REPORT NO.
EPA-600/2-76-008
2.
4. TITLE AND SUBTITLE
SO2 Control Processes for Non-ferrous Smelters
5. REPORT DATE
January 1976
6. PERFORMING ORGANIZATION CODE
Mathews ^ Faust L. Bellegia,
Charles H.Gooding, and George E.Weant
8. PERFORMING ORGANIZATION REPORT NO.
10. PROGRAM ELEMENT N°-1AI3013/015'
tOAP 21ADC-59/21AUY-40/41
II. CONTRACT/GRANT NO.
9. PERFORMING ORflANIZATION NAME AND ADDRESS
Research Triangle Institute
P.O. Box 12194
Research Triangle Park, North Carolina 27709
58-02-1491
12. SPONSORING AGENCY NAME AND ADDRESS
EPA, Office of Research and Development
Industrial Environmental Research Laboratory
Research Triangle Park, NC 27711
13. TYPE OF REPORT AND PERIOD COVERED
~inal: 6/74-6/75
14. SPONSORING AGENCY CODE
15. SUPPLEMENTARY NOTES
is. ABSTRACT The report reviews and evaluates a number of absorption-based SO2 con-
trol systems and the application of these control systems to those U. S. primary
copper smelters which generate weak SO2-containing gas streams. Capital and oper-
ating cost relationships have been developed for each specific process, covering a
range of gas flows and SO2 concentrations. Separate general costs for gas pretreat-
ment and the end-of-the-line SO2 utilization facilities (i.e. , sulfuric acid, elemental
sulfur, and liquid SO2 plants) are also provided. The 13 U.S. primary copper smel-
ters which currently still generate weak SO2 streams have been reviewed with refer-
ence to their current operation and active programs in hand to control or eliminate
weak SO2 streams. Appropriate SO2 control processes have been matched with the
individual smelters and related capital and operating costs have been developed from
the earlier established cost relationships.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b.lDENTIFIERS/OPEN ENDED TERMS
c. COSATI Field/Group
Air Pollution
Absorption
Sulfur Oxides
Smelters
Copper
Sulfuric Acid
Sulfur
Cost Analysis
Air Pollution Control
Stationary Sources
Elemental Sulfur
Liquid SO2
13B
07B
11F
14A
' DISTRIBUTION STATEMENT
Unlimited
19. SECURITY CLASS (This Report}
Unclassified
383
20. SECURITY CLASS (This page)
Unclassified
22. PRICE
:*•>.
A Form 2220-1 (9-73)
371
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