Final Report
Contract CPA 22-69-78
FEASIBILITY STUDY OF NEW SULFUR OXIDE
CONTROL PROCESSES FOR APPLICATION
TO SMELTERS AND POWER PLANTS
Part II: The Wellman-Lord SO2 Recovery
Process for Application to Smelter Gases
Prepared for:
U.S. DEPARTMENT OF HEALTH, EDUCATION, AND WELFARE
NATIONAL AIR POLLUTION CONTROL ADMINISTRATION
DURHAM, NORTH CAROLINA
STANFORD RESEARCH INSTITUTE
Menlo Park, California 94025 • U.S.A.
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STANFORD RESEARCH INSTITUTE
Menlo Park, California 94025 • U.S.A.
Final Report
Contract No. CPA 22-69-78
FEASIBILITY STUDY OF NEW SULFUR OXIDE
CONTROL PROCESSES FOR APPLICATION
TO SMELTERS AND POWER PLANTS
Part II: The Wellman-Lord SO2 Recovery
Process for Application to Smelter Gases
By: KONRAD T. SEMRAU
Prepared for:
U.S. DEPARTMENT OF HEALTH, EDUCATION, AND WELFARE
NATIONAL AIR POLLUTION CONTROL ADMINISTRATION
DURHAM, NORTH CAROLINA
SRI Project PMU-7923
Approved:
N. K. HIESTER, Director
Physical Sciences (Materials}
C. J. COOK, Executive Director
Physical Sciences Division
Copy No.
88
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CONTENTS
FOREWORD vii
I INTRODUCTION 1
II OBJECTIVES 5
III SUMMARY 7
IV PROCEDURES 11
A. Models 11
B. Cost Factors 11
C. Preparation of Technical Data and Cost Estimates ... 11
V PROCESS DESCRIPTION 13
VI PROCESS DATA AND COST ESTIMATES 23
A. Sulfur Oxides Recovery .... 23
B. Gas Cleaning 24
C. Cost Estimates for Wellman-Lord System 24
D. Cost Estimates for Model Control Schemes 26
VII GENERAL DISCUSSION
A. Evaluation of the Wellman-Lord SO Recovery System . . 61
£
B. Disposal of By-Products 63
REFERENCES 65
APPENDIXES
A. SMELTER MODELS A-l
B. FACTORS AND CONDITIONS ASSUMED IN ESTIMATING
CONTROL SYSTEM COSTS B-l
C. GAS CLEANING SYSTEM C-l
D. AUXILIARY CONTACT SULFURIC ACID PLANTS D-l
iii
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APPENDIXES (Continued)
E; SULFUR DIOXIDE REDUCTION PLANTS . E-l
ILLUSTRATIONS
Figure 1 Wellman-Lord SO Recovery System for Smelter Gases ... 16
Figure 2 Capital Costs of Wellman-Lord Absorbers 3O
Figure 3 Annual Costs of Wellman-Lord Absorbers 31
Figure 4 Capital Costs of Wellman-Lord Chemical Recovery Sections 34
Figure 5 Annual Costs of Wellman-Lord Chemical Recovery Sections . 35
Figure 6 Model A Copper Smelters, Control Scheme CA-1. — Effect
on Copper Production Cost ......... 52
Figure 7 Model A Copper Smelters, Control Scheme CA-2 — Effect
on Copper Production Costs 53
Figure 8 Model B Copper Smelters, Control Scheme CB-1 — Effect
on Copper Production Costs 54
Figure 9 Model B Copper Smelters, Control Scheme CB-2 — Effect
on Copper Production Costs 55
Figure 10 Model A Zinc Smelters, Control Scheme ZA-1 — Effect
on Zinc Production Costs 56
Figure 11 Zinc Smelter Models B, C, and D; Control Schemes ZB-1,
ZC-1, and ZD-1 — Effects on Zinc Production Costs ... 57
Figure 12 Zinc Smelter Models B, C, and D; Control Schemes ZB-2,
ZC-2, and ZD-2 — Effects on Zinc Production Costs ... 58
Figure 13 Model A Lead Smelters, Control Scheme LA-1 — Effect
on Lead Production Costs 59
Figure C-l Gas Cleaning System for Smelter Gases C-2
Figure C-2 Capital Cost of Gas Cleaning System for Smelter Gases . . C-4
Figure C-3 Total Annual Cost of Gas Cleaning System for Smelter
Gases C-8
Figure D-l Capital Costs of Auxiliary Sulfuric Acid Plants D-3
iv
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ILLUSTRATIONS (Continued)
Figure D-2 Annual Costs of Auxiliary Sulfuric Acid Plants D-4
Figure E-l Capital Costs of Sulfur Dioxide Reduction Plants .... E-4
Figure E-2 Annual Costs of Sulfur Dioxide Reduction Plants .... E-5
TABLES
Table
Utilities, Labor, and Chemical Requirements of
Wellman-Lord Absorbers
Table
Table
Table
Table
Table
Table
Table
Table
Table
Table
Table
Table
Table
Table
II
III
IV
V
VI
VII
VIII
IX
X
XI
XII
XIII
XIV
XV
Capital and Annual Costs of Wellman-Lord Absorbers
(SO less than 2.5%)
2
27
28
Capital and Annual Costs of Wellman-Lord Absorbers
(SO0 greater than 2.5%) ..... 29
&
Utilities, Labor, and Chemical Requirements of
Wellman-Lord Chemical Recovery Systems 32
Capital and Annual Costs of Wellman-Lord Chemical
Recovery Systems . 33
Gas Flow Rates and Sulfur Dioxide Emissions from
Model Smelters 36
Recovery of Sulfur Dioxide from Smelter Gases by
Wellman-Lord System
Description of Smelter Model Control Schemes
Copper Smelter Model A — Costs of Control Scheme CA-1
Copper Smelter Model A — Costs of Control Scheme CA-2
Copper Smelter Model B — Costs of Control Scheme CB-1
Copper Smelter Model B — Costs of Control Scheme CB-2
Zinc Smelter Model A — Costs of Control Scheme ZA-1
Zinc Smelter Model B — Costs of Control Scheme ZB-1
Zinc Smelter Model B — Costs of Control Scheme ZB-2
37
38
40
41
42
43
44
45
46
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TABLES (Continued)
Zinc Smelter Model C — Costs of Control Scheme CZ-1 .
Zinc Smelter Model C — Costs of Control Scheme CZ-2 .
Zinc Smelter Model C — Costs of Control Scheme CZ-1 .
Zinc Smelter Model D — Costs of Control Scheme ZD-2 .
Lead Smelter Model A — Costs of Control Scheme LA-1 .
47
48
49
50
51
Table XVI
Table XVII
Table XVIII
Table XIX
Table XX
Table A-l Smelter Models — Summary of Gas Compositions and
Flow Rates and of Sulfur Emissions A-2
Table A-2 Metal Production by Model Smelters A-3
Table C-l Gas Cleaning System for Smelter Gases — Liquid Flows
and Power Requirements * C-5
Table C-2 Gas Cleaning System for Smelter Gases — Utilities and
Operating Labor Requirements C-6
Table C-3 Gas Cleaning System for Smelter Gases — Capital and
Total Annual Costs C-7
Table D-l Labor and Utilities Requirements of Auxiliary
Sulfuric Acid Plants D-2
Table D-2 Capital and Annual Costs of Auxiliary Sulfuric
Acid Plants D-3
Table E-l Labor, Utilities, and Raw Material Requirements of
Sulfur Dioxide Reduction Plants K-2
Table E-2 Capital and Annual Costs of Sulfur Dioxide Reduction
Plants E-3
vi
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FOREWORD
The final report for this study is presented in four separate and
independent parts:
Part I: The Monsanto Cat-Ox Process for Application to
Smelter Gases
Part II: The Wellman-Lord SO Recovery Process for Application
to Smelter Gases
Part III: The Monsanto Cat-Ox Process for Application to Power
P^ant Flue Gases
Part IV: The Wellman-Lord SO Recovery Process for Application
to Power Plant Flue Gases
Information for use in this study was supplied to Stanford Research
Institute by Monsanto Company and Wellman-Lord, Inc. under terms of con-
fidentiality agreements between the U.S. Department of Health, Education
and Welfare, Stanford Research Institute, and each of the cooperating
companies. In accordance with the agreements, Monsanto Company and
Wellman-Lord, Inc. have reviewed and released the parts of the report
dealing with their respective processes. The rights of prior review
and release are designed solely to permit the cooperating companies to
assure themselves that no proprietary or confidential data are being
revealed; they are not intended to restrict Stanford Research Institute's
rights and responsibilities to report its conclusions so long as there
is no incidental disclosure of confidential information. Accordingly,
the release of the reports by Monsanto and Wellman-Lord does not imply
that these companies necessarily concur in all or any of the opinions,
judgments, or interpretations of fact expressed by the author, who
assumes sole responsibility for the report content.
vii
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I INTRODUCTION
Under the Systems Study for Control of Emissions — Primary Non-
ferrous Smelting Industry (Contract No. PH 86-68-85), Arthur G. McKee
& Company and its subcontractor, Stanford Research Institute, carried
out evaluations of a number of sulfur oxide control processes as they
might be applied to offgases from nonferrous smelting. To permit eval-
uation of the technical and economic feasibility of these control pro-
cesses, a number of models of smelters were created. Stanford Research
Institute carried out the studies necessary to determine the availability
of markets for sulfur by-products open to smelters in various areas, and
the allowable production costs that the smelters would have to attain in
order to break even on the sulfur recovery operations.
The Division of Process Control Engineering of the National Air
Pollution Control Administration (DPCE-NAPCA) desires to extend the use-
fulness of the foregoing study by adding to it technical and economic
evaluations of new and potentially promising sulfur oxide control pro-
cesses. It also wishes to evaluate the same new processes for application
to power plants. Completion of these preliminary evaluations of the
processes will help determine their potential commercial acceptability.
DPCE-NAPCA has a specific interest in at least two control processes
being offered commercially, the Monsanto Cat-Ox process and the Wellman-
Lord SO Recovery process. However, both processes are proprietary, and,
A
as a matter of policy, DPCE-NAPCA does not wish to obtain proprietary
and confidential information on the processes. It does, nevertheless,
wish to obtain evaluations in nonconfidential terms. Broadly, DPCE-
NAPCA wishes to obtain estimates of the capital and annual costs of the
control systems for each of the assumed applications, together with
appraisals of the technical constraints on each process and of the cur-
rent states of development of the processes.
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Stanford Research Institute was requested by DPCE-NAPCA to carry
out evaluations of the processes under the terms of confidentiality
agreements between the Department of Health, Education, and Welfare,
the owners of the proprietary processes, and Stanford Research Institute.
SRI, acting as a disinterested third party, was to make analyses of the
processes using information obtained from Monsanto Company and Wellman-
Lord, Inc., and to report the results to DPCE-NAPCA without compromising
any of the Monsanto or Wellman-Lord confidential data.
Because of the requirements of confidentiality, it is not permissible
to describe certain features of the Cat-Ox and Wellman-Lord processes.
The corresponding portions of the systems have had to be represented only
in terms of their general functions, and SRI's evaluations of these por-
tions has had to be presented in the form of conclusions without supporting
data or reasoning. In other instances, the parts of the systems could be
described in general, but specific details and design parameters could not
be revealed.
Within the scope of the present project it would obviously have been
impossible to inspect and evaluate independently all the company records
and design data even had the cooperating companies been requested to per-
mit this and had they acceded to the request. The author of this report,
who also conducted the study, evaluated the information provided at his
request, using his own knowledge and relevant data from the literature
and other available sources. Whenever apparent discrepancies or uncer-
tainties were noted in the information, efforts were made to secure veri-
fication or clarification from the companies. In instances where resolution
of questions was not possible, or the information required proved to be
simply unavailable, the author employed his best judgment.
Throughout the following sections of this report, information for
which other sources are not specifically cited was generally obtained
from the cooperating companies and accepted by the author either because
it could be verified from other sources or because it appeared reason-
able. In other instances, information or estimates were provided by
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the companies that could not be verified independently or judged for
reasonableness; in such cases, the companies have been specifically
cited as the sources. In still other instances, the author did not
accept the information or estimates provided and in some cases sub-
stituted his own; such cases have also been specifically noted.
The cooperating companies, at their own option and through sub-
stantial efforts, provided the basic capital cost estimates for the
model control systems and the information for estimation of operating
and maintenance costs. The author in this case acted as a reviewer
rather than as an estimator. The estimates were checked for reason-
ableness and for possible errors or omissions. For some components
and cost factors, the author modified the estimates, or substituted
others of his own where he judged them to be more appropriate than
those supplied to him. The author also prepared cost estimates for
some auxiliary systems, using separate data sources.
By specification of the power plant and smelter models, and by
review of the results, an effort was made to ensure that the cost esti-
mates for both control systems were made on strictly comparable bases.
Although it is unlikely that this objective has been met fully, the
deviations are probably within the precision of the estimates them-
selves .
For convenience, and at the request of DPCE-NAPCA, this final report
is presented in four separate and independent parts. This part deals
only with the Wellman-Lord SO Recovery process as applied to control of
^
sulfur oxides in the offgases from copper, lead, and zinc smelters.
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II OBJECTIVES
The objectives of the part of the study covered by this report are
as follows:
1. To prepare a block flow diagram of the Wellman-Lord system
showing its configuration and its relation to the smelting
processes in which the sulfur oxides are generated.
2. To present estimated mass and volume flow balances for the
Wellman-Lord system.
3. To prepare preliminary engineering estimates of the capital
investment and the total annual cost (including both fixed
and variable charges) for the Wellman-Lord system.
4. To apply the estimates prepared in (3) to the gas streams
of the model smelters created in the previous Systems Study
for Control of Emissions — Primary Nonferrous Smelting
Industry, and to determine the total annual cost for each
model control system. From the estimates of total annual
costs, secondary estimates are to be made of the corresponding
incremental costs of producing the nonferrous metals, both
on the gross basis (without allowance for by-product recovery
credits) and on the net basis (with allowance for by-product
recovery credits).
5. To make a qualitative appraisal of technical constraints on
the application and operation of the control system.
6. To appraise (quantitatively, to the extent permitted by
available data) the economic constraints on the application
of the control system.
7. To assess the current state of development of the Wellman-
Lord system, identifying any technological deficiencies whose
elimination might enhance the applicability of the system to
smelter gases.
The accomplishment of the objectives is subject to any restrictions
that may be imposed under the terms of the confidentiality agreement
between the Government, Wellman-Lord, Inc., and Stanford Research
Institute.
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Ill SUMMARY
The Wellman-Lord SO Recovery process is a cyclic absorption-
£
desorption process for producing concentrated sulfur dioxide from waste
gas streams, such as power plant flue gas and smelter offgases. It
is adaptable to use on gas streams containing sulfur dioxide at both
low and relatively high concentrations.
The basic Wellman-Lord system is made up of two basically different
parts: (1) the absorption section, in which the sulfur dioxide is removed
from the flue gas, and (2) the chemical recovery section, in which the
concentrated sulfur dioxide is recovered from the spent absorbent and the
absorbent is regenerated for return to the absorption section. The ab-
sorbent is a solution of sodium sulfite. In the absorption cycle, the
sulfite ion reacts with sulfur dioxide, forming bisulfite. In the recovery
cycle, the reaction is reversed by application of heat, releasing sulfur
dioxide and regenerating the sulfite.
The absorption step is conventional and has been employed commercially
as well as studied experimentally. The regeneration step apparently repre-
sents an innovation; it is intended to reduce the amount of steam required
where regeneration is accomplished by direct stripping of the bisulfite
solution with steam. The Wellman-Lord regeneration (or chemical recovery)
system appears to be capable of operating with a steam demand in the range
of about one-half to one-third that required in a conventional steam strip-
ping system. The regeneration equipment is relatively expensive, but not
sufficiently so for the associated fixed charges to offset the advantage
of the reduction in steam consumption. Since the regeneration system has
not yet received patent protection, it cannot be described at this time.
The Wellman-Lord system treated in this study is essentially a con-
ceptual design. Components of the system have been tested individually,
but a complete, integrated system has not yet been operated. The regene-
ration system is basically different from that used in another version
7
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of the process using potassium sulfite solution as the absorbent. The
latter system, which was tested in a demonstration plant on power plant
flue gases, was generally unsatisfactory. However, the Wellman-Lord
system as designed for use with the sodium-base absorbent appears to be
technically feasible.
It is essential that the gas entering the Wellman-Lord system be
as free as possible from dust, fume, and vapor or gaseous contaminants,
such as arsenic trioxide, hydrogen chloride, hydrogen fluoride, and
sulfur trioxide. Solid particles in the absorbent solution may produce
a variety of mechanical and chemical problems, while acidic gases and
vapors will consume absorbent. In addition, a number of the contaminants
may tend to catalyze the oxidation of the sulfite to sulfate. It was
therefore specified by SRI as a model condition that the gas should be
assumed to be cleaned to the same degree as it would be for use in a
contact sulfuric acid plant.
In the Wellman-Lord process, as in all similar processes, some
portion of the sulfite (or sulfur dioxide) is oxidized to sulfate, from
which the sulfur dioxide cannot be regenerated. The sulfate must be
purged from the system. The fraction of the sulfur dioxide that is
oxidized has not been established. Wellman-Lord originally estimated
that the amount would be very low, and provided no system for recovering
the sodium base from the purge stream of absorbent solution, which was
expected to be small. The author anticipated that the amount of sulfur
dioxide oxidized would be much larger than that originally projected by
Wellman-Lord, and assumed the amount to be 1.0 percent of that absorbed,
although the actual percentage is unpredictable. He therefore assumed
.that the purged absorbent solution would be recausticized with lime to
recover the sodium base.
The purging of the sulfate results in a loss of sulfur dioxide in
the sodium sulfite present in the purge stream. If as much as 4 to 6
percent of the sulfur dioxide should be oxidized, a substantial fraction
of the total sulfur dioxide will be lost in the purge, and a system for
separating the sodium sulfate and sulfite will be needed.
8
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The product of the Wellman-Lord system, concentrated sulfur dioxide,
has only very limited markets. Hence, it will in general be necessary
either to convert the sulfur dioxide to sulfuric acid in an auxiliary
contact acid plant, or to reduce it to elemental sulfur in an auxiliary
reduction plant. The markets for sulfuric acid are limited in the areas
of the United States where there are the greatest number of smelters with
uncontrolled emissions, and where elemental sulfur is a more desirable
by-product.
In the model studies (using the smelter models created by Arthur
G. McKee & Company), the Wellman-Lord process was applied conceptually
in 12 control schemes producing sulfuric acid or elemental sulfur. Where
sulfuric acid was to be the product, the Wellman-Lord system was applied
only to gas streams containing 2 percent or less of sulfur dioxide, and
conventional contact acid plants were applied to richer gases. Where
elemental sulfur was to be the product, the Wellman-Lord process was
applied to all the gas streams to concentrate the sulfur dioxide for
subsequent reduction.
The limiting economic factor in the application of the Wellman-Lord
process proved to be the capital and operating costs of the chemical
recovery section, which dominated the total cost. The indicated costs
of recovery of sulfur dioxide from lean gases were high, as with all
other recovery processes. With richer gases, the recovery cost remained
high because the size of the chemical recovery section was proportional
to the quantity of sulfur dioxide regardless of the concentration in the
gas treated. For all of the control schemes producing elemental sulfur,
the cost of production exceeded by factors of two or more any realizable
price that might be obtained for the sulfur. Of those control schemes
producing sulfuric acid, only one attained a production cost that might
be reasonably competitive in some smelter areas; in this scheme, sulfur
dioxide was concentrated from reverberatory furnace gases in a copper
smelter and fed to a conventional sulfuric acid plant operating on the
roaster and converter gases.
9
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For the various smelter models and control schemes, the gross
incremental costs of producing the metals (before allowance for by-product
credits) ranged as follows:
Copper: 1.5 to 3.6 cents/lb
Zinc: 1.1 to 2.3 cents/lb
Lead: 0.6 to 1.1 cents/lb
The by-product credits that might be realized would not in most
cases reduce these costs by a large factor.
10
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IV PROCEDURES
A. Models
The models of the hypothetical smelters, which are presented in
Appendix A, are the same as those formulated and used in the previous
1 f*
study of smelter emission control by Arthur G. McKee &. Company. Some
supplemental conditions, necessary to the specific analyses of the
Cat-Ox and Wellman-Lord processes, were specified by SRI and are also
given in Appendix A.
B. Cost Factors
The cost factors used in making the cost estimates for the control
16
systems were also taken from the McKee report, and are presented in
Appendix B. Some supplemental factors needed specifically for the present
study were specified by SRI.
The estimates of the prices that might be obtained for sulfur by-
16
products were also taken from the McKee report.
C. Preparation of Technical Data and Cost Estimates
Wellman-Lord, Inc. prepared capital cost estimates for gas absorbers
of three different capacities and for chemical recovery sections of five
different capacities , and also estimated the utilities and labor require-
ments. The author reviewed these estimates and accepted most of them.
The cost breakdowns for individual components of the system were supplied
by Wellman-Lord for review. Most of the components consisted of conventional
items of chemical process equipment, and the costs were estimated by Wellman-
Lord on the basis of quotations from vendors. The author surveyed these
estimates for general reasonableness and consistency with the model speci-
fications , but did not attempt further verification.
Some of the original estimates of system requirements or costs made
11
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by Wellman-Lord were not accepted, and the author supplied his own
estimates. Such instances are noted in the following sections of this
report. The changes were discussed with Wellman-Lord, and differences
were generally resolved.
The cost estimates for the gas cleaning system and the auxiliary
sulfuric acid and sulfur dioxide reduction plants were prepared by the
author (see Appendixes C, D, and E), from published data.
12
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V PROCESS DESCRIPTION
The Wellman-Lord SO Recovery process is a cyclic absorption-
£
desorption process for producing concentrated sulfur dioxide from
various concentrations in waste gas streams, such as power plant flue
gas and smelter offgases. It is adaptable to use on gas streams con-
taining sulfur dioxide at both low and relatively high concentrations.
Its product, concentrated sulfur dioxide, has very limited markets, but
can be used in sulfuric acid manufacture or be reduced to elemental sul-
fur. Therefore, the Wellman-Lord process potentially has two principal
uses: (1) concentration of sulfur dioxide from dilute gas streams (those
containing less than about 2 to 3 percent of sulfur dioxide) for use in
sulfuric acid plants, and (2) concentration of sulfur dioxide from both
dilute and rich gas streams for subsequent reduction to elemental sulfur.
If sulfuric acid is the desired product, and the gas stream contains at
least 3.5 to 4 percent of sulfur dioxide, it is practical to treat the
6
gas directly in a contact sulfuric acid plant.
The absorption step in the Wellman-Lord process cycle is conventional,
and in closely related forms has been used commercially and been studied
widely on an experimental basis by numerous investigators. Wellman-Lord,
Inc. believes that some elements of the absorber design are novel and is
seeking patent protection for them. Although the writer questions — on
purely technical grounds — whether the design features are actually nove^.,
it is not critical to the present evaluation of the basic process whether
the claimed innovations in equipment are real or not.
The absorbent consists of a solution of sodium or potassium sulfite.
In the absorption cycle, the sulfite ion reacts with sulfur dioxide as
indicated forming bisulfite:
S03~ -i- S02 + H20 = 2 HS03~
In the desorption cycle, the reaction is reversed by application of heat
13
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releasing sulfur dioxide and regenerating the sulfite:
The absorption step has been employed commercially in the pulp and
paper industry (usually with sodium and ammonia as bases) for production
1 5 19
of cooking liquor; ' ' in this case no regeneration step is employed.
The complete cycle, including the desorption step, has been studied
13 14
experimentally ' and employed occasionally on a limited commercial
13
scale. However, the major economic limitation has been the quantity
of steam required for the absorbent regeneration. In the conventional
14
approach to regeneration, the rich absorbent (containing a large frac-
tion of bisulfite) is stripped of sulfur dioxide with steam in a counter-
current tower. At the bottom of the tower, where the lean stripped
absorbent solution emerges, the bisulfite has been largely converted to
sulfite and the back pressure of sulfur dioxide above the solution is
low. The practical limits of regeneration of the absorbent are set by
the required tower height and the quantity of steam.
Various attempts have been made to avoid the steam stripping of the
15
absorbent. Johns tone, for example, proposed introducing a chemical
regeneration cycle. The Wellman-Lord process employs crystallization of
salts from the absorbent as part of a procedure for reducing the steam
requirements for regeneration. Various proposed forms of the process are
3 17 18
described in Wellman-Lord patents. ' ' One, which uses the potass ium-
3
base salts, is described broadly in a Belgian patent, and more specifically
23 , 22
in a paper by Watt. It was studied on a pilot plant scale and was
23
later investigated in a demonstration plant. The anticipated steam eco-
nomy was not attained, and development has been shifted to another form of
the process, which employs the sodium-base salts.
The sodium -base process, which is used as the basis for the present
study, has a regeneration process that is essentially different from that
employed where the potassium-base salts are used. It is the regeneration
process that constitutes the apparently novel feature of the Wellman-Lord
system. Applications have been made for patents, but patent protection
14
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has not yet been received. Consequently, it is not permissible at this
1;ime to describe the basic regeneration process or the system for carrying
it out. The corresponding portions of the regeneration system are there-
fore omitted from the flowsheet, Fig. 1.
The system considered in the present study has not been operated in
a complete, integrated system, but experiments have been carried out
with individual components. It is essentially a conceptual design backed
by relevant experience with the potassium-base pilot plant and demonstra-
tion plant and by data from the tests of system components. A commercial
plant using the process is to go into service in 1970 on the tail gas
from a contact sulfuric acid plant.
The regeneration system devised by Wellman-Lord avoids the direct
steam stripping of the absorbent in a stripping column, and reduces the
steam requirement substantially below that of the stripper. During
regeneration, the amount of sulfur dioxide recovered per unit quantity
of absorbent circulated depends upon the concentration of the solution
and upon the fraction of the salt that is initially bisulfite. Hence,
attaining maximum steam economy (the least consumption of steam per
unit quantity of sulfur dioxide recovered) requires that the rich absor-
bent leaving the absorption process be as concentrated as possible, and
that the conversion of the sulfite to bisulfite also be as nearly complete
as possible.
The upper limit of concentration of the sodium sulfite in the lean
absorbent solution fed to the absorber is set by the solubility of the
salt. As the sulfite is converted to the more soluble bisulfite by
reaction with sulfur dioxide (the actual crystalline compound is the
purosulfite instead of the bisulfite), the solution becomes less nearly
saturated. Hence, a portion of the water in the solution can be evapo-
rated into the gas stream passing through the tower. However, a practical
limit on the amount of water removed in this manner is imposed by the
need to avoid bringing the rich absorbent solution to saturation with
respect to sodium pyrosulfite. If saturation is reached, plugging of
parts of the t-bsorber and absorbent outlet piping by pyrosulfite crystals
15
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ABSORPTION SECTION
GAS TO STACK
1
K
Ul
CO
ac
CLEAN
SMELTER
GAS
110°F
CHEMICAL RECOVERY
SECTION
LEAN
ABSORBENT
RICH
ABSORBENT
sf
a
— *• C(
S02
COMPRESSOR
3NDENSER
I
TANK
tu
1
tu
a
1
so2
HEATER
CONDENSER
1
TANK
1
PUMP
I CONDENSATE
AND SO2
— i
"I
ABSORBENT
REGENERATION
SYSTEM
I I
8
STEAM <
CONDENSATE
TA -7923-10
FIGURE 1 WELLMAN-LORD S02 RECOVERY SYSTEM FOR SMELTER GASES
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may occur.
The circulation rate of the absorbent is set by the quantity of
sulfur dioxide to be removed from the gas stream. If the gas being
treated is relatively rich in sulfur dioxide (as, for instance, a
smelter gas), it is possible in a countercurrent absorption tower to
convert somewhat more of the sulfite to bisulfite than if the gas is
more dilute; a higher back pressure of sulfur dioxide above the solution
leaving the tower is permissible because the partial pressure of sulfur
dioxide in the entering gas stream is also higher. However, the addi-
14
tional fractional conversion of the sulfite is not large, and the gain
in capacity of the absorbent is ignored in conservative design of a
system. The sulfur dioxide capacity is therefore taken as constant
regardless of the composition, of the gas stream treated.
The steam economy of the system can be improved by staging the
regeneration process , but the economic limits of this approach are set
by the concomitant increase in capital investment. The optimum balance
between operating costs and depreciation must therefore be determined.
In the Wellman—Lord process, as in all similar processes employing
the sulfite-bisulfite absorption and recovery cycle (or even absorption
without regeneration), some portion of the sulfite is oxidized to sulfate,
from which the sulfur dioxide cannot be regenerated. The sulfate must
therefore be purged from the system. The fraction of the sulfite that
may be oxidized in the Wellman-Lord system has not yet been established;
it may, in fact, vary widely depending upon such factors as the oxygen
content of the gas stream, the temperature, and the presence of contami-
nants that may act as oxidation catalysts. Wellman-Lord adds hydroquinone
to the absorbent solution as an oxidation inhibitor, but the effective-
ness of this measure is questionable. This point is discussed below.
The basic Wellman-Lord system is made up of two basically different
parts, the absorption section and the chemical recovery section (see
Fig. 1). The absorption section is composed only of the absorber and
the fan. Its size, and hence its cost, is dependent almost solely upon
17
-------
the volume of gas handled, with only E. minor cost component related to
the concentration of the sulfur dioxide in the gas. The size and cost
of the chemical recovery section are dependent upon the total quantity
of sulfur dioxide recovered. Since the quantity of sulfur dioxide ab-
sorbed per unit quantity of absorbent is constant, the total volume of
absorbent circulated is directly proportional to the total quantity of
sulfur dioxide.
Removal of particulate matter from the gas before the latter enters
the absorber is a critical aspect of the operation of the Wellman-Lord
process. (It will evidently be equally critical to any other cyclic
process for concentration of sulfur dioxide.) Particulate matter enter-
ing the absorption system may produce mechanical problems in the chemical
recovery section (e.g., plugging lines and wearing pumps) and may inter-
fere in other ways with the operation of the system. It must be purged
from the system by filtration or centrifuging and may carry with it
occluded absorbent. If the filtered particulate is washed to recover
the absorbent, the resulting dilution of the solution increases the
steam requirements for the regeneration process. The particulate matter
may also include elements that will act as catalysts for the oxidation
of the sulfite.
The gas stream may contain, in addition to sulfur dioxide, other
gases and vapors that are deleterious to the system. Examples are sul-
fur trioxide, hydrogen chloride, hydrogen fluoride, and arsenic trioxide.
A wet scrubbing system is therefore an essential part of the preliminary
gas cleaning system, both for removing soluble gaseous contaminants and
for cooling and humidifying the gas stream. Sulfur trioxide (and sul-
furic acid mist), hydrogen chloride, and hydrogen fluoride all form
sodium salts, consuming the absorbent base and requiring purging from
the system. It was specified by SRI (Appendix A) that the gas entering
the system should be assumed to: (1) have been purified in a gas cleaning
system typical of those used with contact sulfuric acid plants, (2) con-
tain residual contaminants in concentrations at least as low as are
specified by Donovan and Stuber, and (3) be at 110 F and saturated with
18
-------
water vapor. The capital and operating costs of the gas cleaning system
-avre 'eliarguil -against sulfur oxide recovery, but were estimated separately
by SRI from literature data (see Appendix C) and are presented as separate
i terns.
The clean smelter gas is forced up through the countercurrent absorb-
ing tower where it contacts the absorbent solution on a series of trays
and the sulfur dioxide is absorbed. The treated gas leaving the top tray
passes through an entrainment separator where entrained droplets of the
absorbent solution are removed, and then is discharged to the stack.
Arrangements for gas-liquid contacting in the tower vary somewhat depend-
ing upon the concentration of the sulfur dioxide in the gas. For sulfur
dioxide concentrations under 2.5 percent the liquid throughput is rela-
tively low, and special provisions are made to ensure good gas-liquid
contact. These provisions (not indicated in Figure 1) make the capital
cost somewhat higher than that for an absorber for a richer gas, in which
the liquid throughput is higher and no extra measures are necessary to
ensure effective gas-liquid contacting.
A portion of the sulfuric acid mist remaining in the cleaned smelter
gas will be collected in the absorber. The amount is not readily predict-
able, but the author assumed 90 percent.
The gas pressure drop across the absorber is approximately 12 inches
of water. The fan is sized and powered to supply the draft for movement
of the gas through the gas cleaning system, which has an additional pres-
sure drop of 14 inches of water. The fan and large circulating pumps of
the Wellman-Lord system are driven by backpressure steam turbines. The
exhaust steam (at 50 psig) is sufficient to supply the requirements of
the chemical recovery section. Only small pumps and other equipment are
electrically driven.
The rich absorbent solution leaving the bottom of the absorbing
tower is pumped to the chemical recovery section, where the sulfur dioxide
is released from the bisulfite solution and the sulfite is reformed. A
mixture of the sulfur dioxide and steam leaves the system and goes first
o
to the main condenser, which is operated at about 200 F. Most of the
19
-------
steam is condensed, but the temperature is high enough that little of
the sulfur dioxide is absorbed in the condensate, which can be returned
to the system as makeup water. The hot, wet sulfur dioxide stream
leaving the receiving tank of the main condenser enters the second
o
condenser, where it is cooled to about 120 F. Host of the remaining
water vapor condenses, but the condensate contains a high concentration
of dissolved sulfur dioxide. This second condensate stream must there-
fore be fed to the condensate stripper, where the sulfur dioxide is
recovered by stripping with a small amount of steam. The overhead from
the stripper joins the main gas stream from the second condenser receiving
tank. At this point the main stream of sulfur dioxide gas is at about
o
120 F and saturated with water vapor, which then constitutes less than
12 percent by volume of the gas stream. The stripper overhead contains
a much higher concentration of water vapor, but is only a small fraction
of the main gas stream. The wet product sulfur dioxide is heated to
prevent subsequent condensation of water vapor, then compressed to about
8 psig for delivery to an adjacent plant for subsequent manufacture of
sulfuric acid or reduction to elemental sulfur.
To prevent the buildup of an excessive sodium sulfate concentration
in the absorbent, a purge stream of the lean absorbent solution is with-
drawn. In the original design supplied by Wellman-Lord for this study,
no provision was made for recovery of the sodium base from the absorbent
purge stream because it was believed that the amount of sulfite (or sulfur
dioxide) oxidized to sulfate would be very low. The author believes that
the amount of sulfur dioxide oxidized will probably be much higher and
that recovery of the purged base will be an economic necessity. If the
amount of sulfur dioxide oxidized does not exceed perhaps 1 to 2 percent
of that absorbed, it may be feasible to recausticize the sodium sulfate-
sodium sulfite mixture with lime, precipitating the sulfate and sulfite.
The recovered sodium hydroxide solution may have to be carbonated to pre-
cipitate calcium ions remaining in solution. If calcium ions remain in
the solution, they may eventually produce scaling problems in the recovery
system. As a guesstimate, the capitalized cost of operating such a recov-
ery system might be about $12 to $15 per ton of sodium hydroxide recovered,
20
-------
including the cost of lime, makeup sodium hydroxide, operation, and
depreciation of equipment. The higher figure, $15/ton, is assumed in
this study.
In the purging of sulfate, a portion of the sulfur dioxide is lost
because of purging of the associated sulfite. The author calculated the
average sulfur dioxide loss to be about 6.5 percent of that absorbed
from the flue gas, assuming that 1.0 percent is oxidized and that 90 per-
cent of the sulfuric acid mist in the clean smelter gas is collected in
the absorber to form additional sulfate. If the amount of sulfur dioxide
oxidized should be as much as perhaps 4 to 6 percent of that absorbed,
the sodium hydroxide could still be recovered by the recausticization
process, but the amount of sulfur dioxide lost in purging would be a
substantial fraction of the total. In such case, there is a need for a
system for separating and purging the sodium sulfate that will not result
in loss of the sulfite.
There will be a gradual buildup of other contaminants (such as
chlorides, fluorides, and solid particles) in the absorbent, originating
from the residuals in the clean smelter gas. These contaminants will
also have to be purged from the system along with the sulfate.
21
-------
VI PROCESS DATA AND COST ESTIMATES
A. Sulfur Oxides Recovery
In the smelter models (Appendix A), the distribution of the emitted
sulfur oxides between sulfur dioxide and sulfur trioxide is in most cases
not given. Where the concentration of sulfur trioxide in the gas is not
specified, the concentration of sulfur dioxide is calculated on the
assumption that all of the sulfur is in the form of sulfur dioxide (see
Table A-l). In fact, a variable quantity of sulfur trioxide (equivalent
to a few percent of the total sulfur emission) will be present in all
cases, and will be removed in the gas cleaning system that precedes the
sulfur dioxide recovery system. However, for the sake of simplification,
it was assumed in this study that all of the sulfur was in the form of
sulfur dioxide and was treated in the sulfur dioxide control system.
The inputs and outputs of sulfur from the recovery systems were calculated
on this basis.
Tfi
In the McKee study, different collection efficiencies for sulfur
dioxide were assumed for the various control systems modeled, as appeared
appropriate for the systems considered. In the present study, a uniform
standard of performance was set for both of the recovery systems (Cat-Ox
and Wellman-Lord) that were treated (see Appendix A). Both systems appeared
to be capable of attaining or exceeding the efficiencies specified, at least
within the ranges of input sulfur dioxide concentrations for which their
application was appropriate. The calculated collections of sulfur dioxide
were based on attainment of the specified efficiencies.
The overall efficiency of sulfur dioxide recovery with the Wellman-Lord
system is reduced by the inefficiencies of the auxiliary sulfuric acid or
sulfur dioxide reduction plant (see Appendixes D and E). The losses from
the auxiliary plant could be reduced by applying a Wellman-Lord absorber
to the tail gas and returning the spent absorbent to the same chemical
23
-------
recovery section that serves the main gas absorption system. However,
no such additional absorbers were included in the control system models.
B. Gas Cleaning
The costs (capital and operating) of preliminary gas cleaning (see
Appendix C) were taken to be the same for both the Cat-Ox and Wellman-Lord
systems. However, the costs for the gas cleaning and sulfur dioxide recov-
ery systems are presented separately in this report, and are combined only
in final totals.
The capital cost for the gas cleaning system does not include that
for the fan, which is assigned to the sulfur dioxide recovery system.
However, the operating cost for the gas cleaning system does include the
cost of electrical power consumed by the fan that moves the gas through
the gas cleaning system.
The maximum allowable concentrations of impurities in the cleaned gas
(Appendix A and Reference 7) are apparently higher, in at least a few
instances, than the levels that are favored by some other authorities or
that are attained in some installations. For sulfuric acid mist, Donovan
7
and Stuber give an upper limit of 0.022 grain/std cu ft, which is assumed
4
in this study, whereas Carter favors a maximum of 0.0012 grain/cu ft.
12
Heinrich and Anderson report attainment of residual acid mist and arse-
nic concentrations as low as O.OOOOO22 grain/cu ft with a wet electrostatic
precipitator.
C. Cost Estimates for Wellman-Lord System
To permit relatively rapid estimation of costs for the various model
control schemes, curves of capital and annual costs were prepared for the
Wellman-Lord system (Figs. 2 through 5). Separate curves were derived for
the absorption section (based on gas flow) and for the chemical recovery
section (based on sulfur dioxide input). Similar curves were prepared for
the gas cleaning system (Appendix C), the auxiliary sulfuric acid plant
(Appendix D), and the sulfur dioxide reduction plant (Appendix E). Costs
24
-------
for the various system components were then read off from the appropriate
curves to derive cost estimates for the control schemes.
The capital cost for the absorption section (Fig. 2) was taken from
Wellman-Lord estimates, except that the author added an allowance for
ductwork connecting the outlet of the absorber to the stack that was
assumed to be an existing part of the smelter. The capital cost for the
chemical recovery section (Fig. 4) was taken from the Wellman-Lord esti-
mates without change. Utilities and labor requirements for the absorption
section (Table I) and chemical recovery section (Table IV) were based on
Wellman-Lord estimates. The ratios of sodium hydroxide and hydroquinone
makeup to sodium sulfate purged were also supplied by Wellman-Lord, but
the author made his own estimates of the amount of sulfate formed (see
Section V). The original estimate of sodium hydroxide entrainment loss
from the absorber was tenfold higher than that adopted by the author,
who believes that the values given in Table I should be reasonable unless
reentrainment should take place in the eliminator.
Losses of sodium hydroxide due to entrainment and reaction with sul-
furic acid mist are proportional to gas flow and were charged against the
absorption section. Those due to oxidation of the sulfur dioxide are pro-
portional to the quantity of sulfur dioxide and were charged to the chem-
ical recovery section, even though the oxidation actually takes place in
the absorber (except for possible autoxidation of the sulfite).
All of the steam used to drive the turbine-powered fans is later
used for solvent regeneration in the chemical recovery section, and the
entire steam cost was charged to the latter (Table IV). Since the steam
condensate is returned to the boiler plant, the charge for the steam is
reduced by a credit for the value of the condensate.
The absorber sections for use on the more dilute gases require
electrical drives for equipment not employed where richer gases are
treated (Table I). The chemical recovery sections (particularly the
smaller ones) for use on plants treating dilute gases have a relatively
high proportion of electrical drives (Table IV). Where richer gases are
25
-------
being treated and steam requirements for regeneration are higher, it is
practical to drive large rotating equipment in the chemical recovery
section with turbines, since the exhaust steam can be used for regenera-
tion.
The computations of annual costs for the absorption and chemical
recovery sections are presented in Tables II, III, and V. The effect of
the sulfur dioxide concentration in the gas on the annual cost of the
chemical recovery section was so small that only a single cost curve was
drawn in Fig. 5.
D. Cost Estimates for Model Control Schemes
The gas flow rates and sulfur dioxide emissions from the model
smelters are summarized in Table VI. Because the copper converter gas
flow rate varies widely, the rate used in designing gas handling equip-
ment was assumed to be higher than the average (see Appendix A). The
performances of the WeiIman-Lord systems on the various smelter gas
streams are presented in Table VII.
The smelter model control schemes treated are described in Table
VIII, and the cost estimates are given in Tables IX through XX. All the
capital and annual costs were taken from the curves in this report except
those for the conventional contact sulfuric acid plants used in Control
Schemes CA-1 and CB-1, which were taken from curves in Reference 16.
Conversion in these acid plants was assumed to be 97%.
In Figs. 6 through 13 the changes in the unit production costs of
the metals, resulting from the application of the control schemes to the
model smelters, are shown graphically as functions of the net selling
price of the sulfur by-product at the smelter. (The quantities of metal
produced by the model smelters are given in Appendix A.) The net selling
price is equal to the gross selling price less sales cost and overhead,
and will, of course, be dependent upon the particular location of the
plant.
26
-------
Table I
UTILITIES, LABOR, AND CHEMICAL REQUIREMENTS
OF WELLMAN-LORD ABSORBERS
Gas Flow Rate
ACFM
(110°F, 1 atm, sat 'd)
37,000
229,000
580,000
Cost
Factor
Electric
Power
(kw)
6.6
19.8
52.6
1.0£ /kwh
Sodium Hydroxide
Makeup
(Entrainment Loss)
(Ib/hr)
0.37
2.29
5.80
$65.00/ton2
Sodium Hydroxide
Makeup
(Acid Mist Reaction)
(Ib/hr)
24.3
150
381
$15.00/ton3
Hydroquinone
(Ib/hr)
0.05
0.27
0.70
82£/lb
Labor
/man-hr\
\ day J
12
$3.75/hr
to
1 Only with S02 under 2.5%.
2 Purchased sodium hydroxide.
O
Recovered sodium hydroxide.
-------
Table II
CAPITAL AND ANNUAL COSTS OF WELLMAN-LORD ABSORBERS
Less than 2.5%)
Item
Capital Cost
Annual Cost
A. Depreciation
B. Direct Operating Cost
1. Labor
2. Supervision
3. Payroll benefits
4. Maintenance materials
5. Factory supplies
6. Electricity
7. Chemicals
C. Indirect Costs
1. Controllable
2. Noncon troll able
Total Annual Cost
Cost
Basis
$
$/yr
$/yr
Gas Flow Rate-ACFM
(110°F, 1 atm, sat'd)
37,000
196,000
13,070
14,850
—
3,710
5,880
980
520
1,860
10,370
5,880
57,120
229,000
600,000
40,000
14,850
—
3,710
18,000
3,000
1,570
11,250
16,430
18,OOO
126,810
580, 000
1,060,000
70,670
14,850
—
3,710
31,800
5,300
4,170
28,670
23,330
31,800
214,300
28
-------
Table III
CAPITAL AND ANNUAL COSTS OF WELLMAN-LORD ABSORBERS
(SQ3 greater than 2.5%)
Item
Capital Cost
Annual Cost
A. Depreciation
B. Direct Operating Cost
1. Labor
2. Supervision
3. Payroll benefits
4. Maintenance materials
5. Factory supplies
6. Electricity
7. Chemicals
C. Indirect Costs
1. Controllable
2. Noncon troll able
Total Annual Cost
Cost
Basis
$
$/yr
$/yr
Gas Flow Rate-ACFM
(110°F, 1 atm, sat'd)
37 , 000
186,000
12,400
14,850
—
3,710
5,580
930
—
1,860
10,220
5,580
55,130
229,000
580,000
38,670
14,850
—
3,710
17 , 400
2,900
—
11,250
16,130
17 , 400
122,310
580,000
1,030,000
68,670
14,850
—
3,710
30,900
5,150
—
28,670
22,880
30,900
205,730
29
-------
10
e
C
o
8
o
I
u
1.0
0.1
10.000
i i r
TT
i L 1 I 1 I I I I
I I I i I I I I
100,000
GAS FLOW RATE—cu ft/min (110°F. 1 atm. safd)
FIGURE 2 CAPITAL COSTS OF WELLMAN-LORD ABSORBERS
1.000,000
TA-7923-11
30
-------
1000
o
-a
•o
I
o
te
O
o
D
Z
<
o
100
10
10,000
11
100,000
GAS FLOW RATE—cu ft/min (110°F, 1 atm, safd)
FIGURE 3 ANNUAL COSTS OF WELLMAN-LORD ABSORBERS
1.000,000
TA-7923-12
31
-------
Table IV
UTILITIES, LABOR, AND CHEMICAL REQUIREMENTS OF
WELLMAN-LORD CHEMICAL RECOVERY SYSTEMS
Sulfur
Dioxide
Input
(Ib/hr)
6,000
12,000
24,000
35,000
49,000
Cost
Factor
Steam
< Ib/hr)
35,100
70,200
140,400
204,800
286,700
75^/1000 Ib
Electric Power
(kw)
SO2>
1;5%
104
190
309
438
557
802=
0.5%
318
348
349
438
~ 557
IcAwh
Process
Water
(Ib/hr)1
3,510
7,020
14,000
20,500
28,700
5^/1000 Ib
Cooling
Water
Makeup
(gpm)
40
80
160
240
330
2C/1000 gal
Sodium
Hydroxide
< Ib/hr)2
466
932
1,870
2,720
3,810
$15/ton3
Hydroquinone
(Ib/hr)
0.84
1.68
3.37
4.90
6.86
82CVlb
Operating
Labor
/man-hr\
\ day /
56
$3.75/hr
Supervision
/man-hr\
\ day /
12
$4.75Ar
w
to
*• Steam condensate.
2 Baaed on oxidation of 1.0 percent of input sulfur dioxide.
3 Recovered sodium hydroxide.
4 Two operators per shift plus one day laborer.
-------
Table V
CAPITAL AND ANNUAL COSTS OF
WELLMAN-LORD CHEMICAL RECOVERY SYSTEMS
CO
CO
Item
Capital Cost
Annual Cost
A. Depreciation
B. Direct Operating Cost
1. Labor
2. Supervision
3, Payroll benefits
4. Maintenance materials
5. Factory supplies
6. Steam
7. Electricity
(a) S02 = 0.5%
(b) S02 > 1.5%
8. Process water
9. Cooling water makeup
10. Chemicals
C. Indirect Costs
1. Controllable
2. Noncontrollable
Total Annual Cost
(a) S02 = 0.5%
(b) S02 > 1.5%
Cost
Basis
$
$/yr
$/yr
Sulfur Dioxide Input — Ib/hr
6,000
1,590,000
106 , 000
69,300
18,810
22,030
47,700
7,950
208,490
25,190
8,240
1,390
380
33,130
67,900
47,700
655,970
639,020
12,000
2,290,000
152,670
69,300
18,810
22,030
68,700
11,450
416,990
27,560
15,050
2,780
760
66,270
78,400
68,700
1,004,420
991,910
24,000
3,300,000
220,000
69,300
18,810
22,030
99 , 000
16,500
833,980
27,640
24,470
5,540
1,520
132,960
93,560
99,000
1,639,840
1,636,670
35,000
4,000,000
266,670
69,300
18,810
22,030
120,000
20,000
1,216,500
34,690
34,690
8,120
2,280
193,390
104,060
120,000
2,195,850
2,195,850
49,000
4,800,000
320,000
69,300
18,810
22,030
144,000
24,000
1,703,000
44, 110
44,110
11,370
3,140
270,860
116,060
144,000
2,890,680
2,890,680
-------
10
i—r
TTT
1—I I I I I JJ
•5
|
i
8
u
o
0.1
1000
J_ L 1 I I I I I I
I
I I I I I I I
10,000
SULFUR DIOXIDE INPUT—Ib/hr
100,000
TA-7923-13
FIGURE 4 CAPITAL COSTS OF WELLMAN-LORD CHEMICAL RECOVERY SECTIONS
34
-------
10
o
•o
c
o
8
o
D
Z
O
1.0
0.1
1.000
i i r
1 I i i i m
I
10.000
SULFUR DIOXIDE INPUT—Ib/hr
100,000
TA-7923-14
FIGURE 5 ANNUAL COSTS OF WELLMAN-LORD CHEMICAL RECOVERY SECTIONS
35
-------
Table VI
GAS FLOW RATES AND SULFUR DIOXIDE EMISSIONS
FROM MODEL SMELTERS
Smelter
Copper
Zinc
Lead
Model
A
B
A
B
C
D
A
Process Unit
Reverberatory Furnace
Converters
1
Roaster
Reverberatory Furnace
Converters
1
Converters + Roaster
2
Roaster
Roaster
Sinter Plant
Roaster
Sinter-Roaster
Sinter Plant
Sulfur
Dioxide
Concentration
(%)
1.89
3.82
11.2
0.91
3.78
6.05
7.1
0.9
0.5
0.8
2.0
5.0
Gas Flow Rate
CFM
(110°F, 1 atm, sat'd)
Small
81,000
75,100
102 , 900
29,800
48,800
67 , 300
87,500
97,100
117,200
22,400
154,400
55,500
12,400
Medium
162 , 000
150,200
193,300
59,600
97,700
134,600
165,600
194,200
225,200
44,800
409,600
166,900
24,700
Large
243,000
225,300
292,900
89,400
146,600
201,800
248,200
291,200
337,600
67,100
49,500
Sulfur Dioxide Emission
(Ib/hr)
Small
14,100
26,480
30,630
4,100
23,600
54,230
14,930
12,850
2,570
5,720
Medium
28,200
52,970
61,270
8,200
47,200
108,470
29,870
30,300
30,870
11,430
Large
42,300
79,450
91,900
12 , 300
70,800
162,700
44,800
22,870
CO
05
1 Design basis for converters.
2 Design basis for converters + roaster.
-------
Table VII
RECOVERY OF SULFUR DIOXIDE FROM SMELTER
GASES BY WELLMAN-LORD SYSTEM
Smelter
Copper
Zinc
Lead
Model
A
B
A
B
C
D
A
Process Unit
Reverberatory Furnace
Converters
Roaster
Reverberatory Furnace
Converters
Converters + Roaster
Roaster
Roaster
Sinter Plant
Roaster
Sinter— Roaster
Sinter Plant
Sulfur Dioxide
Absorption
Efficiency
(%)
95
98
98
94.51
98
98
98
94.45
90
93.75
95
98
Sulfur Dioxide Absorbed
(Ib/hr)
Small
13,400
25,950
30,020
3,870
23,130
53,150
14,630
12,140
2,310
5,600
Medium
26,790
51,910
60,040
7,750
46,260
106,300
29,270
28,400
29,300
11,200
Large
40,190
77,860
90,060
11,600
69,380
159,500
43,900
22,400
Sulfur Dioxide Recovered
(Ib/hr)1
Small
12,500
24,300
28,100
3,620
21,600
49,700
13,700
11,400
2,160
5,240
Medium
25,000
48,500
56,100
7,250
43,300
99,400
27,400
26,600
27,400
10,500
Large
37,600
72,800
84,200
10,800
64,900
149,000
41,000
20,900
CO
Loss of sulfur dioxide assumed to be 6.5 percent of amount absorbed.
-------
Table VIII
DESCRIPTION OF SMELTER MODEL CONTROL SCHEMES
Smel ter
Moael
Control
Scheme
No.
Overall SO,
£
Recovery
Description
Copper
A
CA-1
95.4
CA-2
90.3
Contact sulfuric acid plant on
converter gas.
Wellman-Lord recovery plant on
reverberatory furnace gas. Con-
centrated SO3 to contact acid
plant on converter gas.
Individual Wellman-Lord absorbers
on converter gas and reverberatory
furnace gas.
Single Wellman-Lord chemical re-
covery section. Concentrated SO2
to SO reduction plant.
3
Copper
B
CB-1
96.6
CB-2
91.0
Contact sulfuric acid plant on
mixed roaster and converter gas.
Wellman-Lord recovery plant on re-
verberatory furnace gas. Concen-
trated SO,, to contact acid plant
2
on roaster and converter gas.
Individual Wellman-Lord absorbers
on reverberatory furnace gas and
on mixed roaster and converter gas
Single Wellman-Lord chemical re-
covery section. Concentrated SO2
to SO- reduction plant.
O
Zinc
A
ZA-1
91.3
Wellman-Lord recovery plant on roaster
gas
plant.
Concentrated SO3 to SO2 reduction
(continued)
38
-------
Table VIII (concluded)
Smelter
Model
Control
Scheme
No.
Overall S0r
<
Recovery
Description
Zinc
B
ZB-1
ZB-2
92.0
87.3
1. Individual Wellman-Lord absorbers
on roaster gas and sinter plant
gas.
2. Single Wellman-Lord chemical re-
covery section. Concentrated SO^
to contact sulfuric acid plant.
1. Individual Wellman-Lord absorbers
on roaster gas and sinter plant
gas.
2. Single Wellman-Lord chemical re-
covery section. Concentrated SO3
to SO reduction plant.
Zinc
C
ZC-1
ZC-2
92.0
87.3
Wellman-Lord recovery plant on roaster
gas. Concentrated SO to contact sul-
furic acid plant.
Wellman-Lord recovery plant on roaster
gas. Concentrated SOg to SOfe reduc-
tion plant.
Zinc
D
ZD-1
ZD-2
93.2
88.5
Wellman-Lord recovery plant on sinter-
roaster gas. Concentrated SQg to con-
tact sulfuric acid plant.
Wellman-Lord recovery plant on sinter-
roaster gas. Concentrated SOg to SQg
reduction plant.
Lead
A
LA-1
91.3
Wellman-Lord recovery plant on sinter
plant gas. Concentrated SO
reduction plant.
2toS02
39
-------
Table IX
COPPER SMELTER MODEL A—COSTS OF CONTROL SCHEME CA-1
Item
Capital Investment
1. Gas Cleaning System
2. Wei loan -Lord System
(a) Absorption
(b) Chemical recovery
3. Contact Acid Plant
Total
Annual Cost
1. Gas Cleaning System
2. Wellman-Lord System
(a) Absorption
(b) Chemical recovery
3. Contact Acid Plant
Total
Sulf uric Acid Produced
(100% basis)
Gross Costs Before Credits
1. Production Cost of Acid
2. Incremental Cost of
Producing Metal
Cost or Credit
Small Plant
$ 1,830,000
319,000
2,420,000
5,500,000
$10,069,000
$ 358 ,OOO
75,000
1,090,000
1,400,000
$ 2,923,000
229,200 tons/yr
$12.75/ton
1.93?/lb
Medium Plant
$ 2,950,000
490,000
3,500,000
11,000,000
$17,940,000
$ 583 , 000
105 , 000
1,760,000
2,800,000
$ 5,248,000
458,400 tons/yr
$11.45/ton
1.73
-------
Table X
COPPER SMELTER MODEL A— COSTS OF CONTROL SCHEME CA-2
Item
Capital Investment
1. Gas Cleaning System
(a) Reverberatory furnace
(b) Converters
2. Wei Iraan -Lord Absorbers
(a) Reverberatory furnace
(b) Converters
3. Wei Iraan -Lord Chemical Recovery
4. SO2 Reduction Plant
Total
Annual Cost
1. Gas Cleaning System
(a) Reverberatory furnace
(b) Converters
2. Wellman-Lord Absorbers
(a) Reverberatory furnace
(b) Converters
3. Wellman-Lord Chemical Recovery
4. SO2 Reduction Plant
Total
Elemental Sulfur Produced
Gross Costs Before Credits
1. Production Cost of Sulfur
2. Incremental Cost of
Producing Metal
Cost or Credit
Small Plant
$ 1,830,000
2,160,000
319 , OOO
355,000
4,260,000
1,560,000
$10,484,000
$ 358,000
423,000
75,000
81,000
2,380,000
730,000
$ 4,047,000
67,500 short tons/yr
$59.96/short ton
2.67C/lb
Medium Plant
$ 2,950,000
3,330,000
490,000
525,000
8,520,000
2,340,000
$18,155,000
$ 583 , 000
665 , 000
105,000
112 , 000
4,760,000
1,230,000
$ 7,455,000
134,800 short tons/yr
$55.30/short ton
2.45?/lb
Large Plant
$ 3,900,000
4,400,000
630,000
675,000
12,780,000
2,950,000
$25,335,000
$ 780,000
900,000
132,000
140,000
7,140,000
1,700,000
$10,792,000
202,500 short tons/yr
$53.29/short ton
2.37C/lb
41
-------
Table XI
COPPER SHELTER MODEL B--COSTS OF CONTROL SCHEME CB-1
Item
Capital Investment
1. Gas Cleaning System
2. Wellman-Lord System
(a) Absorption
(b) Chemical recovery
3. Contact Acid Plant
Total
Annual Cost
1. Gas Cleaning System
2. Wellman-Lord System
(a) Absorption
(b) Chemical recovery
3. Contact Acid Plant
Total
Suit uric Acid Produced
(100% basis)
Gross Costs Before Credits
1. Production Cost of Acid
2. Incremental Cost of
Producing Metal
Cost or Credit
Small Plant
$1,280,000
233,000
1,260,000
7,100,000
$9,873,000
$ 256,000
62,000
500,000
1,700,000
$2,518,000
339,900 tons/yr
$7.41/ton
1.66£/lb
Medium Plant
$ 2,O8O,OOO
358,000
1,820,000
14,200,000
$18,458,000
$ 408,000
82,000
765,000
3,400,000
$ 4,655,000
679,800 tons/yr
$6.85/ton
1.53£/lto
Large Plant
$ 2,750,000
460,000
2,250,000
21,300,000
$26,760,000
$ 540,OOO
100,000
985,000
5,100,000
$ 6,725,000
1,020,000 tons/yr
$6.59/ton
1.47?/lb
42
-------
Table XII
COPPER SMELTER MODEL B—COSTS OF CONTROL SCHEME CB-2
Item
Capital Investment
1. Gas Cleaning System
(a) Reverberatory furnace
(b) Roaster and converters
2 . Wellman— Lord Absorbers
(a) Reverberatory furnace
(b) Roaster and converters
3. Wellman-Lord Chemical Recovery
4. 303 Reduction Plant
Total
Annual Cost
1. Gas Cleaning System
(a) Reverberatory furnace
(b) Roaster and converters
2. Wellman-Lord Absorbers
(a) Reverberatory furnace
(b) Roaster and converters
3. Wellman-Lord Chemical Recovery
4. SO2 Reduction Plant
Total
Elemental Sulfur Produced
Gross Costs Before Credits
1. Production Cost of Sulfur
2. Incremental Cost of
Producing Metal
Small Plant
S 1,280,000
2,350,000
233,000
380,000
7,200,000
1,930,000
§13,373,000
§ 256,000
460,000
62 , 000
86 , 000
3,680,000
960,000
$ 5,504,000
97,800 short tons/yr
$56.28/short ton
3.63?/lb
Cost or Credit
Medium Plant
$ 2,080,000
3,690,000
358,000
575,000
12,600,000
2 , 9OO , OOO
$22,203,000
$ 408,000
740,000
82 , 000
121,000
6,900,000
1,650,000
S 9,901,000
195,600 short tons/yr
$50.62/short ton
3.26C/lb
Large Plant
S 2,750,000
4,850,000
460,000
740,000
17,800,000
3,650,000
330,250,000
S 540,000
1 , 000 , 000
100,000
152,000
10,200,000
2,300,000
$14,292,000
293,100 short tons/yr
S48.76/short ton
3.13
-------
Table XIII
ZINC SMELTER MODEL A
COSTS OF CONTROL SCHEME ZA-1
Item
Capital Investment
1. Gas Cleaning System
2. Wellman-Lord System
(a) Absorption
(b) Chemical recovery
3. SO2 Reduction Plant
Total
Annual Cost
1. Gas Cleaning System
2. Wellman-Lord System
(a) Absorption
(b) Chemical recovery
3. S02 Reduction Plant
Total
Elemental Sulfur Produced
Gross Costs Before Credits
1. Production Cost of Sulfur
2. Incremental Cost of
Producing Metal
Cost or Credit
Small Plant
$ 760,000
137,000
2,530,000
870, 000
$4,297,000
$ 159,000
49,000
1,150,000
400,000
$1,758,000
25 , 100 short tons/yr
$70.04/short ton
1.55
-------
Table XIV
ZINC SMELTER MODEL B
COSTS OF CONTROL SCHEME ZB-1
Item
Capital Investment
1. Gas Cleaning Systems
(a) Roaster
(b) Sinter plant
2. Wellman-Lord Absorbers
(a) Roaster
(b) Sinter plant
3. Wellman-Lord Chemical Recovery
4. Contact Acid Plant
Total
Annual Cost
1. Gas Cleaning Systems
(a) Roaster
(b) Sinter plant
2. Wellman-Lord Absorbers
(a) Roaster
(b) Sinter plant
3. Wellman-Lord Chemical Recovery
4. Contact Acid Plant
Total
Sulfuric Acid Produced (100% basis)
Gross Costs Before Credits
1. Production Cost of Acid
2. Incremental Cost of Producing
Metal
Cost or Credit
$2,850,000
1,410,000
472 , 000
252,000
2,520,000
645 , 000
$8,149,000
$ 560,000
278,000
102,000
65,000
1,140,000
212,000
$2 , 357 , 000
80,550 tons/yr
$29.26/ton
2 . 08£/lb
45
-------
Table XV
ZINC SMELTER MODEL B
COSTS OF CONTROL SCHEME ZB-2
Item
Cost or Credit
Capital Investment
1. Gas Cleaning System
(a) Roaster
(b) Sinter plant
2. Wellman-Lord Absorbers
(a) Roaster
(b) Sinter plant
3. Wellman-Lord Chemical Recovery
4. SOa Reduction Plant
Total
$2,850,000
1,410,000
472,000
252,000
2,520,000
860,000
$8,364,000
Annual Cost
1. Gas Cleaning System
(a) Roaster
(b) Sinter plant
2. Wellman-Lord Absorbers
(a) Roaster
(b) Sinter plant
3. Wellman-Lord Chemical Recovery
4. SOa Reduction Plant
Total
$ 560,000
278,000
102,000
65,000
1,140,000
400,000
$2,545,000
Elemental Sulfur Produced
24,900 short tons/yr
Gross Costs Before Credits
1. Production Cost of Sulfur
2. Incremental Cost of Producing
Metal
$102/short ton
2.25^/lb
46
-------
Table XVI
ZINC SMELTER MODEL C
COSTS OF CONTROL SCHEME ZC-1
Item
Cost or Credit
Capital Investment
1. Gas Cleaning System
2. Wellman-Lord System
(a) Absorption
(b) Chemical recovery
3. Contact Acid Plant
Total
$ 5,550,000
860,000
3,600,000
990,000
$11,000,000
Annual Cost
1. Gas Cleaning System
2, Wellman—Lord System
(a) Absorption
(b) Chemical recovery
3. Contact Acid Plant
Total
$ 1,150,000
176,000
1,840,000
304,000
$ 3,470,000
Sulfuric Acid Produced
(100% basis)
158,000 tons/yr
Gross Costs Before Credits
1. Production Cost of Acid
2. Incremental Cost of
Producing Metal
$21.96/ton
47
-------
Table XVII
ZINC SHELTER MODEL C
COSTS OF CONTROL SCHEME ZC-2
Item
Cost or Credit
Capital Investment
1. Gas Cleaning System
2. Wellman-Lord System
(a) Absorption
(b) Chemical recovery
3. SOa Reduction Plant
Total
$ 5,550,000
860,000
3,600,000
1,270,000
$11,280,000
Annual Cost
1. Gas Cleaning System
2. Wellman-Lord System
(a) Absorption
(b) Chemical recovery
3. SOs Reduction Plant
Total
$ 1,150,000
176,000
1,840,000
585,000
$ 3,751,000
Elemental Sulfur Produced
48,800 short tons/yr
Gross Costs Before Credits
1. Production Cost of Sulfur
2. Incremental Cost of
Producing Metal
$76.86/short ton
48
-------
Table XVIII
ZINC SMELTER MODEL D
COSTS OF CONTROL SCHEME ZD-1
Item
Cost or Credit
Capital Investment
1. Gas Cleaning System
2. Wellman-Lord System
(a) Absorption
(b) Chemical recovery
3. Contact Acid Plant
Total
$3,000,000
500,000
3,650,000
1,010,000
$8,160,000
Annual Cost
1. Gas Cleaning System
2. Wellman-Lord System
(a) Absorption
(b) Chemical recovery
3. Contact Acid Plant
Total
$ 595,000
107,000
1,870,000
310,000
$2,882,OOO
Sulfuric Acid Produced
(100% basis)
162,800 tons/yr
Gross Costs Before Credits
1. Production Cost of Acid
2. Incremental Cost of
Producing Metal
$17.70/ton
1.28£/lb
49
-------
Table XIX
ZINC SMELTER MODEL D
COSTS OF CONTROL SCHEME ZD-2
Item
Cost or Credit
Capital Investment
1. Gas Cleaning System
2. Wellman-Lord System
(a) Absorption
(b) Chemical recovery
3. SOa Reduction Plant
Total
$3,000,000
500,000
3,650,000
1,310,000
$8,460,000
Annual Cost
1. Gas Cleaning System
2. Wellman-Lord System
(a) Absorption
(b) Chemical recovery
3. SOg Reduction Plant
Total
$ 595,000
107,000
1,870,000
600,000
$3,172,000
Elemental Sulfur Produced
50,300 short tons/yr
Gross Costs Before Credits
1. Production Cost of Sulfur
2. Incremental Cost of
Producing Metal
$63.06/short ton
50
-------
Table XX
LEAD SMELTER MODEL A—COSTS OF CONTROL SCHEME LA-1
Item
Capital Investment
1. Gas Cleaning System
2. Wellman-Lord System
(a) Absorption
(b) Chemical recovery
3. SO2 Reduction Plant
Total
Annual Cost
1. Gas Cleaning System
2. Wellman-Lord System
(a) Absorption
(b) Chemical recovery
3. SO2 Reduction Plant
Total
Elemental Sulfur Produced
Gross Costs Before Credits
1. Production Cost of Sulfur
2. Incremental Cost of
Producing Metal
Cost or Credit
Small Plant
$ 500,000
95,000
1,530,000
500,000
$2,625,000
$ 115,000
44 , 000
625,000
260,000
$1,044,000
9,610 short tons/yr
$lO9/short ton
l.llc/lb
Medium Plant
§ 810,000
147,000
2,200,000
740,000
$3,897,000
$ 168 , OOO
50,000
960,000
350,000
§1,528,000
19,300 short tons/yr
$79.17/short ton
0.81c/lb
Large Plant
$1,310,000
225,000
3,180,000
1,120,000
$5,835,000
$ 259,000
60 , 000
1,530,000
528,000
$2,377,000
38,300 short tons/yr
$62. 06 /short ton
0.63C/lb
51
-------
o
JO
"5
§
I
u
o
ce
a.
-1
uj
O
i -2
SMALL
MEDIUM
LARGE
I
I
0 5 10 15 20
NET SELLING PRICE Of SULFURIC ACID AT SMELTER — dollars/ton of 100% acid
TA-7923-15
FIGURE 6 MODEL A COPPER SMELTERS: CONTROL SCHEME CA-1—EFFECT
ON COPPER PRODUCTION COSTS
52
-------
o
£
8
o
§
o
o
oc
Ou
«
-2
-3
SMALL
MEDIUM
LARGE
0 20 40 60 80
NET SELLING PRICE OF SULFUR AT SMELTER — dollars/short ton
TA-7923-16
FIGURE 7 MODEL A COPPER SMELTERS: CONTROL SCHEME CA-2-
ON COPPER PRODUCTION COSTS
-EFFECT
53
-------
8
•8
JO
i
o
oc
a.
ui
O
-1
o _2
-3
SMALL
MEDIUM
LARGE
0 5 10
NET SELLING PRICE OF SULFURIC ACID AT SMELTER
15 20
— dollars/ton of 100% acid
TA-7923-17
FIGURE 8 MODEL B COPPER SMELTERS: CONTROL SCHEME CB-1—EFFECT
ON COPPER PRODUCTION COSTS
54
-------
0 20 40 60 80
NET SELLING PRICE OF SULFUR AT SMELTER — dollars/short ton
TA-7923-18
FIGURE 9 MODEL B COPPER SMELTERS: CONTROL SCHEME CB-2—EFFECT
ON COPPER PRODUCTION COSTS
55
-------
N
•5
a
!
g
o
i
g
1-.
-------
0 10 20 30 40
NET SELLING PRICE OF SULFURIC ACID AT SMELTER — dollars/ton of 100% acid
TA-7923-20
FIGURE 11 ZINC SMELTER MODELS B, C, AND D: CONTROL SCHEMES ZB-1,
ZC-1, AND ZD-1—EFFECTS ON ZINC PRODUCTION COSTS
57
-------
o
ft
8
o
z o
o
o
E
0.
UJ
(9
X
-1
-2
-3
ZB-2
ZD-2
120
0 20 40 60 80 100
NET SELLING PRICE OF SULFUR AT SMELTER — dollars/short ton
TA-7923-21
FIGURE 12 ZINC SMELTER MODELS B, C. AND D: CONTROL SCHEMES ZB-2. ZC-2.
AND ZD-2—EFFECTS ON ZINC PRODUCTION COSTS
58
-------
o
a
8
0
o
o
8 -1
o
(C
a.
O
-3
SMALL
MEDIUM
LARGE'
FIGURE 13
20 40 60 80 100 120
NET SELLING PRICE OF SULFUR AT SMELTER — dollars/short ton
TA-7923-22
MODEL A LEAD SMELTERS: CONTROL SCHEME LA-1—EFFECT
ON LEAD PRODUCTION COSTS
59
-------
VII GENERAL DISCUSSION
A. Evaluation of the Wellman-Lord S00 Recovery System
£ —
The Wellman-Lord process using sodium-base absorbent, described
above, appears to be technically feasible. The regeneration of the
absorbent, which is, in the writer's opinion, the only really novel
aspect of the process, appears to be feasible and to offer a signifi-
cant reduction of the steam consumption from the levels required by
direct steam stripping of the rich absorbent. The capital investment
in the Wellman-Lord chemical recovery system is probably much higher
than that for a simple steam-stripping system. However, the associated
fixed charges on the higher investment should be more than compensated
by the reduction of twofold or threefold in the steam requirement.
A major unresolved question concerning the Wellman-Lord system is
the amount of oxidation of the sulfite to sulfate. Some data should be
acquired during 1970 from the commercial plant being constructed for
use on sulfuric acid plant tail gas, but they are unlikely to be gene-
rally applicable even when they do become available. The general
effectiveness of oxidation inhibitors has not been established. There
appears to be no way short of actual experience under the proposed
operating conditions by which the amount of sulfite oxidation can be
predicted with much confidence. Experience with other recovery sys-
259 15 19
terns ' ' ' ' suggests that the amount of the sulfur dioxide oxi-
dized might be in the range of 0.2 to 2.0 percent or even higher,
dependent upon the conditions of the particular case.
Two of the most critical factors in the oxidation of the sulfur
2 19
dioxide are the oxygen concentration in the gas ' and the presence of
oxidation catalysts in the absorbent solution. ' ' In smelter
gases, high oxygen concentrations are frequently present as the result
of air infiltration. The high efficiency cleaning of the smelter gas
ahead of the absorber will be helpful because it will reduce the
61
-------
introduction of potential oxidation catalysts into the absorbent.
However, some of the catalysts are effective at very low concentrations. '
As is discussed above in Section V, oxidation of as little as 4 to
6 percent of the sulfur dioxide can result in loss of a substantial frac-
tion of the total sulfur dioxide during purging of the sulfate. Therefore,
there is potentially a need for an alternative method for separating the
sodium sulfate from the sodium sulfite so that most of the latter can be
returned as such to the absorption cycle.
The type of absorber proposed by WeiIman-Lord is neither the only,
nor necessarily the best, that might be employed. Stone & Webster Engi-
20
neering Corp. prepared estimates of the costs of various types of
absorbers for use with the same sodium sulfite-bisulfite system on appli-
cations to power plant flue gases. Except for the factor of cost, however,
Stone & Webster was unable to present any information that provides a clear
choice between absorber designs.
For smelter applications, the cost of the Wellman-Lord system can be
divided into three parts corresponding to the gas cleaning section, the
absorption section, and the chemical recovery section. (The costs of the
gas cleaning section would be common to any other recovery system.) The
costs of the gas cleaning section and the absorption section are related
to the volume of gas, but the cost of the chemical recovery section is
related only to the quantity of sulfur dioxide handled. The costs of the
auxiliary conversion plants (contact acid or sulfur dioxide reduction) are,
of course, related to the quantity of sulfur dioxide; these costs, too,
would be common to other sulfur-dioxide concentration processes.
The estimated costs for the various model control schemes show clearly
the trends in cost distribution that are related to sulfur dioxide concen-
tration in the gas. For sulfur dioxide concentrations below 1.0 percent,
the capital cost of the gas cleaning section is the largest item, with that
of the chemical recovery section being second and that of the absorption
section much lower. In annual cost, however, the chemical recovery section
is the largest item. With a sulfur dioxide concentration of about 5 percent
62
-------
or higher, the chemical recovery section is by far the largest item in
both capital and annual cost. As the sulfur dioxide concentration is
increased above the range of 5 to 7 percent, the proportionate reduc-
tion in cost (for a given quantity of sulfur dioxide handled) will be
relatively small, because the costs of the gas handling sections are
already a minor component, and the cost of the chemical recovery section
remains the same.
This trend in costs is inherent in the use of sodium sulfite (or
sulfites of other strong bases) as an absorbent. The equilibrium par-
tial pressure of sulfur dioxide is favorable to absorption of sulfur
dioxide from lean gases. However, the capacity for sulfur dioxide is
fixed by the stoichiometric quantity of reactant, and the quantity of
absorbent used has to remain essentially proportional to the total
quantity of sulfur dioxide regardless of concentration. The Wellman-
Lord system — as well as others using the same class of absorbents —
is primarily suited for use on gases containing up to about 3 percent,
or perhaps 4 percent, of sulfur dioxide. For treatment of richer gases,
a more favorable absorbent is one in which the sulfur dioxide is phy-
sically absorbed, or present only in a relatively weak chemical combina-
tion, and the capacity of the solvent is markedly a function of the
partial pressure of the sulfur dioxide. Dimethylaniline is an example
of a solvent for which the sulfur dioxide capacity continues to increase
as the sulfur dioxide concentration in the gas rises to at least 8 or 9
9
percent. With such a solvent, the quantity that must be circulated and
the quantity of steam required for regeneration, per unit quantity of
sulfur dioxide recovered, will continue to decrease as increasingly rich
gas is treated, making the recovery system more economic.
B. Disposal of By-Products
The sizes and availabilities of markets for sulfur by-products
16
producible at smelters have been presented in the McKee report, in
which estimated by-product prices were based on an assumed Gulf Coast
16
sulfur price of $30/long ton, f.o.b. At the time of the previous study,
63
-------
it was estimated that the price of $30 was a likely average for the
period up to 1975. However, sulfur supplies have since shifted from
shortages to surpluses in a period of only about a year, and sulfur
prices have become chaotic. The development of sulfur surpluses has
followed a continuing period of low activity in the market for ferti-
lizers, which provide the largest outlet for sulfur.
In the next five years the withdrawal of marginal producers and a
possible revival of the fertilizer market may cause sulfur prices to
rise from their present lows. However, the figure of $30/1ong ton now
appears to be an optimistic one. It perhaps represents the upper end
of the probable range of Gulf Coast prices to be encountered in the
period to 1975. It cannot even be assumed that the Gulf Coast sulfur
price will continue to maintain its previous status as the base line for
world sulfur prices. Nevertheless, if the above limitations are recog-
nized, the assumption of the $30/long ton price is probably as good as
any that can be made at this time. The estimates of sulfur by-product
prices in smelter areas that are presented in Reference 16 can therefore
be used as a first approximation if it is understood that they probably
represent the most favorable situation that can be anticipated.
Of those model control schemes that produce sulfuric acid, only
Control Scheme CB-1 gives an estimated acid production cost low enough
possibly to be competitive in those areas of the United States where
there are a large number of uncontrolled smelter emissions: (1) Idaho
and Montana, and (2) Arizona, New Mexico, and western Texas. Estimated
acid production costs for all the other control schemes are clearly too
high to be competitive.
Elemental sulfur is a far preferable product, because it can be
shipped much greater distances for sale than can acid, and it can be
easily stored. However, the estimated production costs of sulfur from
all of the control schemes are far higher than the prices that were
generally obtainable even during the sulfur shortage that has Just
passed.
64
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REFERENCES
1. Aho, William A., The Jenssen Exhaust Scrubber — An Effective Air
Protection System, Tappi 52_ (4), 620-623 (Apr. 1969)
2. Applebey, M. P., The Recovery of Sulfur from .Smelter Gases, J.
Soc. Chem. Ind. 56_, 139T-146T (May 1937)
3. Beckwell Process Corporation, Process for Recovery of Sulfur
Dioxide, Belgian Pat. No. 706,449 (May 13, 1968)
4. Carter, B. M. , Sulfuric Acid, Kirk-Othmer Encyclopedia of Chemical
Technology, 1st Ed., 13_, 458-501 (1954)
5. Clement, J. L., and W. L. Sage, Ammonia-Base Liquor Burning and
Sulfur Dioxide Recovery, Tappi 52_ (8), 1449-1456 (Aug. 1969)
6. Connor, J. M., Economics of Sulfuric Acid Manufacture, Chem. Eng.
Progr. 64 (11), 59-65 (Nov. 1968)
7. Donovan, J. R., and P. J. Stuber, Sulfuric Acid Production from
Ore Roaster Gases, J. Metals If) (11), 45-50 (Nov. 1967)
8. Duecker, W. W., and J. K. West (Eds.), "The Manufacture of Sul-
furic Acid," Reinhold Publishing Co., New York (1959)
9. Fleming, E. P., and T. C. Fitt, Liquid Sulfur Dioxide from Waste
Smelter Gases. Use of Dimethylaniline as Absorbent, Ind. Eng.
Chem. 42_ (11), 2253-2258 (Nov. 1950)
10. Fuller, E. C., and R. H. Crist, The Rate of Oxidation of Sulfite
Ions by Oxygen, J. Am. Chem. Soc. 63 (6), 1644-1650 (June 1941)
11. Harris, I. J., and G. H. Roper, The Absorption of Oxygen by
Sodium Sulfite on a Sieve Plate, Can. J. Chem. Eng. 42_ (1), 34-37
(Feb. 1964)
12. Heinrich, R. F., and J. R. Anderson, Electro-Precipitation, in
Cremer and Davies, "Chemical Engineering Practice," Vol. 3, pp.
484-534, Butterworths, London (1956)
13. Johnstone, H. F., Recovery of Sulfur Dioxide from Dilute Gases,
Pulp Paper Mag. Can. 53 (4), 105-112 (Mar. 1952)
65
-------
14. Johnstone, H. F., H. J. Read, and H. C. Blankmeyer, Recovery of
Sulfur Dioxide from Waste Gases. Equilibrium Vapor Pressure over
Sulfite-^isulfite Solutions, Ind. Eng. Chem. 3£ (1), 101-109
(Jan. 1938)
15. Johnstone, H. F., and A. D. Singh, Recovery of Sulfur Dioxide from
Waste Gases. Regeneration of the Absorbent by Treatment with Zinc
Oxide, Ind. Eng. Chem. 32_ (8), 1037-1049 (Aug. 1940)
16. McKee fc Company, Arthur G., Systems Study for Control of Emissions
— Primary Nonferrous Smelting Industry, Final Report to National
Air Pollution Control Administration, June 1969, Contract No. PH 86-
68-85
17. Miller, Leo A., and Jack D. Terrana (to Wellman-Lord, Inc.), Process
for Recovering Sulfur Dioxide from Flue Gas, U.S. Pat. No. 3,477,815
(Nov. 11, 1969)
18. Miller, Leo A. , and Jack D. Terrana (to Wellman-Lord, Inc.), Process
for Recovering Sulfur Dioxide from Gases Containing Same, U.S. Patent
No. 3,485,581 (Dec. 23, 1969)
19. Palmrose, G. V., and J. H. Hull, Pilot Plant Recovery of Heat and
Sulphur from Spent Ammonia -Base Sulphite Pulping Liquor, Tappi 35^
(5), 193-198 (May 1952)
20. Stone fe Webster Engineering Corp., Sulfur Dioxide Scrubbers, Stone
& Webster/Ionics Process, Final Report to National Air Pollution
Control Administration, January 1970, Contract No. CPA 22-69-80
21. Srivastava, R. D., A. F. McMillan, and I. J. Harris, The Kinetics
of Oxidation of Sodium Sulfite, Can. J. Chem. Eng. 4£ (3), 181-184
(June 1968)
22. Terrana, J. D. , and L. A. Miller, Process for Recovery of Sulfur
Dioxide from Stack Gases, Proc. Amer. Power Conf. 30, 627-632 (1968)
(Pub. 1969)
23. Watt, Stuart G., Wellman-Lord S02 Recovery Process, Wellman-Lord,
Ind., Lakeland, Florida (1969)
24. Yodis, A. W. , Applicability of Reduction to Sulfur Techniques to
the Development of New Processes for Removing SO from Flue Gases —
Phase I. Allied Chemical Corp., Interim Report to National Air
Pollution Control Administration for Period June 1, 1968-July 31,
1969 (Draft Copy, Sept. 26, 1969), Contract No. PH 22-68-24.
66
-------
Appendix A
SMELTER MODELS
The models of hypothetical copper, lead, and zinc smelters were
created by Arthur G. McKee & Company, and are described in full in the
A2
McKee report. The essential features of the models directly pertinent
to the present study are summarized here. Some additional model condi-
tions, specific to the present study, were set by SRI, as indicated
below.
In Table A-l are presented the data on gas stream compositions, gas
flow rates, and sulfur emissions for each of the model smelters. The
data on metal production at the model smelters are given in Table A-2.
The remainder of the model conditions are summarized below. Conditions
specified by SRI are noted by an asterisk.
A. Plant Operating Time
330 days/year
24 hours/day
B, Gas Cleaning
1. Gases are assumed to have been cleaned in a hot electro-
static precipitator to a residual dust and fume content of
0.1 grain/std cu ft. The cost of hot, dry gas cleaning is
not charged against sulfur oxides recovery.
*2. The gas leaving the hot precipitator is assumed to be cleaned
in a system typical of those usedfor cleaning feed gases
to a contact sulfuric acid plant (see Appendix C). The
costs (capital and operating) of this secondary cleaning
system are charged against sulfur oxides recovery. The
cleaned gas is assumed to be at 110 F and essentially 1 atm
pressure, and to be saturated with water vapor. It is
further assumed that the residual concentrations of impuri-
ties are at least as low as those Donovan and Stuber^ have
specified to be acceptable in contact sulfuric acid plants:
A-l
-------
Table A-l
SMELTER MODELS
SUMMARY OF GAS COMPOSITIONS AND FLOW RATES AND OF SULFUR EMISSIONS
Plant
Copper
Zinc
Lead
Model
A
B
A
B
C
D
A
Process Unit
Reverberatory
Furnace
Converters
Roaster
Reverberatory
Furnace
Converters
Roaster
Sinter plant
Roaster
Sinter plant
Roaster
Sinter -Boaster
Sinter plant
Gas
Temp.
(°¥)
550
645
550
550
637
600
500
600
300
400
300
400
Gas Composition
(%)
S°2
1.89
3.82
8.00
0.91
3.78
7.1
O.O48
0.9
0.5
0.8
2.0
5.0
S°3
0.1
0.0016
0.29
°2
6.48
5.71
16.1
10.9
18.0
18.0
18.0
16.0
16.0
12.0
H2°
11.7
34.8
0.05
3.2
Gas Flow Rate
1,000 SCFM
(32°F, 1 atm)
Small
69.9
64.8 *
36.0
42.15
58. 043
19.31
29.8
133.2
47.9
—
—
10.67
Medium
139.8
129.6 2
72.0
84.3
116.1 4
38.62
59.6
—
—
353.4
144.0
21.34
Large
209.7
194.4 2
108
126.5
174.1 4
57.93
89.4
—
—
—
—
42.68
Sulfur Equivalent
(short tons/day)
Small
84.6
158.9
183.8
24.6
141.6
89.6
0.9
77.1
15.4
—
—
34.3
Medium
169.2
317.8
367.6
49.2
283.2
179.2
1.8
—
—
181.8
185.2
68.6
Large
253.8
476.7
551.4
73.8
424.8
268.8
2.7
—
—
—
—
137.2
I
to
1 Size recovery plant for 137% of average flow.
2 Size recovery plant for 130% of average flow.
3 Size recovery plant for 130% of average flow if gases combined with roaster gases.
4 Size recovery plant for 123% of average flow if gases combined with roaster gases.
-------
Table A-2
METAL PRODUCTION BY MODEL SMELTERS
Metal Production
Metal
Copper
Size
B
Zinc
Lead
B
C
D
A
Small
Med iutn
Large
Small
Medium
Large
Small
Medium
Large
Small
Medium
Medium
Small
Medium
Large
tons/day
230
460
690
230
460
690
171.4
342.8
514.2
171.4
342.8
342.8
142.9
285.8
571.6
tons/year
75,900
152,000
228,000
75,900
152,000
228,000
56,600
113,000
170,000
56,600
113,000
113,000
47,200
94,300
189,000
A-3
-------
Concentration,
grains/std cu ft,
Contaminant dry basis
Chlorides (as Cl) 0.0005
Fluorides (as F) 0.0001
Arsenic (as As O ) 0.0005
Lead (as Pb) 0.0005
Mercury (as Hg) 0.0001
Selenium (as Se) 0.022
H2S°4 miSt (100%) 0.022
Total Solids (dust) O.O005
*C. Sulfur Dioxide Collection Efficiency
The performance of the primary sulfur dioxide collection system is
to be as follows, defined in terms of collection efficiency or of sulfur
dioxide concentration in the exit gas:
SO Concentration Efficiency (%) or
in Feed Gas Concentration in
(%_) Exit Gas (ppm)
Under 1.0 500 ppm
1.0 to 2.0 95%
Over 2.0 98%
In cases where sulfur dioxide is concentrated for conversion to
another product in a secondary process, the efficiency of the conversion
step is to be taken as a reasonable one for the process assumed.
*D. Cooling Water
The capital and operating costs of the sulfur dioxide control system
are to include the costs of a water cooling tower. Cooling water costs
will be for makeup only.
References
Al. Donovan, J. R., and P. J. Stuber, Sulfuric Acid Production from
Ore Roaster Gases, J. Metals 19_ (11), 45-50 (Nov. 1967)
A2. McKee & Company, Arthur, G., Systems Study for Control of Emissions
— Primary Nonferrous Smelting Industry, Final Report to National
Air Pollution Control Administration, June 1969, Contract No.
PH 86-68-85.
A-4
-------
Appendix B
FACTORS AND CONDITIONS ASSUMED IN
ESTIMATING CONTROL SYSTEM COSTS
The following conditions and cost factors were taken from the
Bl
report by Arthur G. McKee & Company, except that the items marked
with an asterisk are additional conditions specified by SRI.
A. Capital Costs
1. Capital costs do not include:
(a) Working capital
(b) Contingencies
(c) Cost of land
(d) Inventory
(e) Interest on investment
(f) Offsite utilities
(g) Steam generators
(h) Plant access
(i) Cost of dry, hot gas cleaning
2. The cost of the wet cleaning system for the gas is charged to
sulfur oxides control, but is estimated as a separate item
(see Appendix C).
B. Depreciation
15-year, straight-line.
C. Direct Operating Costs (except utilities)
1. Labor $3.75/man-hour
2. Supervision 4.75/man-hour
3. Payroll benefits 25% of labor + supervision
4. Maintenance materials 3% of capital cost of plant
5. Factory supplies O.5% of capital cost of plant
B-l
-------
D. Indirect Costs
1. Controllable indirect costs (maintenance labor, laboratory,
overhead, and supervision)
50% of labor + supervision + maintenance materials
2. Noncontrollable indirect costs (local taxes and insurance)
3% of capital cost of plant
E. Utilities
1. Electrical power l£/kwh
2. Steam 80£/1000 Ib *(4OO psig, 75O°F)
3. natural gas 40£/million Btu
4. Water
(a) Makeup cooling water 2£/1000 gal
(b) Process water 20£/10OO gal
*(c) Boiler feed water 409Y1000 gal
*(d) Retreatment of pro- 38^/1000 gal
cess steam condensates
for boiler use
*(e) Cooling water temper- Q
ature 80 F
F. Waste Acid Disposal
(Weak, waste acid from preliminary gas cooling and cleaning.)
Cost of weak acid disposal 50^/1000 gal
G. Plant Operating Time
330 days/year
24 hours/day
Reference
Bl. McKee & Company, Arthur G., Systems Study for Control of Emissions —
Primary Nonferrous Smelting Industry, Final Report to National Air
Pollution Control Administration, June 1969, Contract No. PH 86-68-85.
B-2
-------
Appendix C
GAS CLEANING SYSTEM
The flowsheet for the gas cleaning system (Fig. C-l) was adapted
from that portrayed for contact sulfuric acid plants in the Arthur G.
C2
McKee & Company final report. The inlet gas conditions were also
assumed to be the same as those specified in the McKee report, except
that a single, average value was taken for the SO concentration (0.2%
o
by volume). Under the conditions assumed, the water vapor condensed
from the incoming gas is sufficient to balance the outflow of weak acid
waste, so that no makeup scrubbing water is required. The cooling water
used in the heat exchanger to cool the recirculated scrubbing water is
assumed to be itself cooled in a water cooling tower. The water cooling
tower is not shown in the flowsheet but is assumed to be part of the sys-
tem. It is treated as a separate unit for convenience in cost accounting,
even though the same water cooling tower would probably in most cases
handle the cooling water for both the gas cleaning and the sulfur dioxide
recovery systems.
The capital cost for the gas cleaning system (not including water
Cl
cooling tower) was taken from the paper by J. M, Connor. Data points
were taken from cost curve No. 2 of Connor's Fig. 5, corresponding to
indirect cooling with a 30 F approach between the cold gas and cooling
water temperatures. The system capacity was converted from the original
basis (tons per day of 100% H SO with gas containing 8% SO ) to the
basis of volumetric gas flow at system exit conditions (110 F, 1 atm, sat'd),
Connor's estimates were presumably made for the type of gas cleaning
system illustrated in his paper, a scrubber-cooler followed by an indirect
tubular type of gas cooler, with an electrostatic mist precipitator for
final cleanup. Although the scrubbing and cooling equipment may therefore
be different from that assumed in the present case, Connor's cost estimates
(which were admittedly approximate) were taken to be sufficiently precise
C-l
-------
Partially Clmnad
GH from Hot
Electrostatic
Precipitate*
OB 600°F
WB 150°F
180°F
O
N
Sludge
•nd
acid vvwte
ELECTROSTATIC
MIST
PRECIPITATOR
GAS SCRUBBER
HUMIDIFIER
TA-7923-7
FIGURE C-1 GAS CLEANING SYSTEM FOR SMELTER GASES
-------
for use in this analysis.
Connor's curves did not extend beyond a capacity equivalent to
about 100,000 CFM of gas, but the plotted curve (Fig. C-2) was extrap-
olated arbitrarily to 1,000,000 CFM. The actual upper limit of validity
probably corresponds to no more than 300 to 400,000 CFM. Above this
size range, the use of multiple units of equipment may make the cost
more nearly proportional to the first power of gas flow rate. However,
none of the individual gas flows in the smelter models exceeds about
400,000 CFM, above which the curve in Fig. C-2 is shown as a broken line.
c
To obtain the curve of capital costs (Fig. C-2) for use in the pres-
ent study, the capital cost of a water cooling tower was added to the
cost of the remainder of the system as adapted from Connor's data.
The gas and water flows through the system were determined by making
heat and material balances based on the conditions shown in Fig. C-l.
Power requirements were estimated from literature data in a few instances,
and calculated from guesstimates of gas pressure drop and pump heads in
the rest. Utilities, labor requirements, and cost factors employed are
summarized in Tables C-l and C-2. The standard cost factors used were the
same as those employed elsewhere in the study (Appendix B).
The total annual costs were calculated for five gas flow capacities
ranging from 10,000 to 1,000,000 CFM (see Table C-3), using the data of
Fig. C-2 and Tables C-l and C-2, and were used to prepare the curve of
annual cost, Fig. C-3. The curve probably is not valid for capacities
greater than 300 to 400,000 CFM, and is shown as a broken line above the
latter level.
References
Cl. Connor, J. M. , Economics of Sulfuric Acid Manufacture, Chem. Eng.
Progr. 64_ (11), 59-65 (Nov. 1968)
C2. McKee & Company, Arthur G., Systems Study for Control of Emissions —
Primary Nonferrous Smelting Industry, Final Report to National Air
Pollution Control Administration, June 1969, Contract No. PH 86-68-85.
C-3
-------
o
I
I I I I I 11
<
10.000
GAS FLOW RATE—cu ft/min (110°F, 1 atm, safd)
FIGURE C-2 CAPITAL COST OF GAS CLEANING SYSTEM FOR SMELTER GASES
C-4
-------
Table C-l
GAS CLEANING SYSTEM FOR SMELTER GAS
LIQUID FLOWS AND POWER REQUIREMENTS
Item
Gas Flow through System
Ap = 14 in. w.c.
Scrubber-Humidifier
Circulated water
Gas Cooling Tower
Circulated water
Cooling Water Circulation
Waste Acid
(21% H2S04)
Water Cooling Tower
Circulated water
Makeup water
Subtotal
Electrostatic Precipitator
Total
Basis
Gas Flow = 1000 CFM at 110°F, 1 a tin, sat 'd
Liquid Flow
7.86 gpm
35.7 gpm
41.7 gpm
0.212 gpm
41.7 gpm
4.17 gpm
Power
3.50 hp
0.43 hp
1,50 hp
1.40 hp
0.06 hp
1.91 hp
8.80 hp
6.57 kw
1.50 kw
8.07 kw
C-5
-------
Table C-2
GAS CLEANING SYSTEM FOR SMELTER GASES
UTILITIES AND OPERATING LABOR REQUIREMENTS
Item
Electrical Power
Gas flow only
Total
Makeup Cooling
Water
Waste Acid
Treatment
Operating Labor
Supervision
Payroll Benefits
Unit Cost
$0.01/kwh
$0.02/1000 gal
$0.50/1000 gal
$3.75/hr
$4.75/hr
25% of Labor +
Supervision
Utilities Basis:
Gas Flow = 1000 CFM @
110°F, 1 atm, sat'd
Quantity
2.62 kw
8.07 kw
4.17 gpm
0.212 gpm
12 hrs/day
2 hrs/day
Annual Cost
($)
207.50
639.14
39.63
50.37
14.8501
3.1351
4.4961
Plant operating time = 330 days = 7920 hrs/yr.
Labor costs are total for plant regardless of size.
C-6
-------
Table C-3
GAS CLEANING SYSTEM FOR SMELTER GASES
CAPITAL AND TOTAL ANNUAL COSTS
Item
Capital Investment
Annual Cost
A. Depreciation
B. Direct Operating Cost
1 . Labor
2. Supervision
3. Payroll benefits
4. Maintenance Materials
5. Factory supplies
6. Electricity
7. Makeup water
8. Waste acid treatment
C. Indirect Costs
1. Controllable
2. Noncontrollable
Total Annual Cost
Cost
Basis
$
$/CFM
$/yr
$/yr
$/yr per
CFM
Gas Flow Rate-CFM at 110°F, 1 atm, saturated
10,000
436,000
43.60
29,100
14,850
3,135
4,496
13,080
2,180
6,390
396
504
15,530
13,080
102,741
10.27
30,000
930,000
31.00
62 , 000
14,850
3,135
4,496
27,900
4,650
19,170
1,190
1,510
22,940
27,900
189,741
6.33
100,000
2,104,000
21.04
140,200
14,850
3,135
4,496
63,120
10,520
63,910
3,960
5,040
40,550
63,120
412,901
4.13
300,000 .
4,480,000
14.93
298,670
14,850
3,135
4,496
134,400
22,400
191,740
11,890
15,110
76,190
134,400
907,281
3.02
1,000,000
10,156,000
10.16
677,100
14,850
3,135
4,496
304,680
50,780
639,140
39,630
50,370
161,330
304,680
2,250,191
2.25
o
I
-------
I I I I Mill
I I I TM 4-
1
t-
<
z
<
UO
O.I
10.000
I Mill
I i 1 M 1 1
100.000
GAS FLOW RATE—cu ft/min (110°F, 1 atm, safd)
1.000.000
TA-7923-9
FIGURE C-3 TOTAL ANNUAL COST OF GAS CLEANING SYSTEM FOR SMELTER GASES
C-8
-------
Appendix D
AUXILIARY CONTACT SULFURIC ACID PLANTS
Estimates of the capital costs of contact sulfuric acid plants for
conversion of concentrated sulfur dioxide to sulfuric acid were derived
Dl *
from Fig. 4 of a paper by Connor. This figure gives the cost of the
acid-making section, only, of a conventional contact plant. This cost
should correspond closely to that of the auxiliary plant, which has no
gas-production section. Sufficient air was assumed to be added to the
D2
sulfur dioxide to give a mole ratio of oxygen to sulfur dioxide of 1.3.
The resulting concentration of sulfur dioxide in the mixed feed gas was
13.9 percent. Conversion of the sulfur dioxide was assumed to be 98 per-
cent.
Estimates of labor and utility requirements (Table D-l) were based
generally on selected values from the references cited (Dl and D2). The
costs for water were modified from those presented in Appendix B. The
process water was assumed to be steam condensate or its equivalent. The
cost for cooling water was taken to include capitalization of cooling
towers, since a cooling tower was not included in the acid plant cost
estimate.
The capital costs for the acid plants are presented in Fig. D-l.
The estimated annual costs are calculated in Table D-2 and presented
graphically in Fig. D-2.
*Note:
As published, Fig. 4 of Connor's paper is presented erroneously.
The data given here were taken from the correct form of Fig. 4, which
has not been published.
D-l
-------
REFERENCES
Dl. Connor, J. M. , Economics of Sulfuric Acid Manufacture, Chem. Eng.
Progra. 6£ (11), 59-65 (Nov. 1968)
D2. Duecker, W. W., and J. K. West (Eds.), "The Manufacture of
Sulfuric Acid," Reinhold Publishing Co., New York (1959)
Table D-l
LABOR AND UTILITIES REQUIREMENTS OF AUXILIARY SULFURIC ACID PLANTS
Item
Electricity
Cooling Water
Process Water
Labor
Supervision
Quantity
per ton
100% H SO
3 4-
produced
25 kwh
7000 gal
20 gal
24 man-hr/day
6 man-hr/day
Unit Cost
1.0? Awh
4
-------
Table D-2
CAPITAL AND ANNUAL COSTS OF AUXILIARY SULFURIC ACID PLANTS
Item
Capital Investment
Annual Cost
A. Depreciation
B. Direct Operating Cost
1 . Labor
2. Supervision
3. Payroll benefits
4. Maintenance materials
5. Factory supplies
6. Electricity
7 . Cooling water
8. Process water
C. Indirect Costs
1. Controllable
2 . Noncontrollable
Total Annual Cost
Cost
Basis
$
$/yr
$/yr
Sulfuric Acid Produced -tons /day
150
450,000
30,000
29,700
9,410
9,780
13,500
2,250
12,380
13,86O
400
26,300
13,500
161,080
500
975,000
65 , 000
29,700
9,410
9,780
29,250
4,880
41,250
46,200
1,320
34,180
29,250
300,220
1,200
1,700,000
113,330
29,700
9,410
9,780
51,000
8,500
99,000
110,900
3,170
45,060
51,000
530,850
D-3
-------
10
•5
I
I 1*
u
0.1
100
I I I I I I I
I I I I I I 11
I I I I I I I I
1000
PLANT CAPACITY — tons/day of 100%
10.000
TA-7923-26
FIGURE D-1 CAPITAL COSTS OF AUXILIARY SULFURIC ACID PLANTS
D-4
-------
10
Tl
o
•o
c
o
1.0
8
o
z
<
O
0.1
I ' I 1 I I 111
100
1000
PLANT CAPACITY — tons/day of 100% H2SO4
FIGURE D-2 ANNUAL COSTS OF AUXILIARY SULFURIC ACID PLANTS
10,000
TA-7923-25
D-5
-------
Appendix E
SULFUR DIOXIDE REDUCTION PLANTS
Estimates of the capital and operating costs of sulfur dioxide
reduction plants were derived from data developed by Allied Chemical
El
Corporation for a conceptual design based on the Asarco process.
Allied presented costs for plants of two sizes, but the estimates for
one were factored from detailed estimates made for the larger of the
two. In Fig. E-l the curve of capital cost was extrapolated (dashed
line) to both smaller and larger sizes of plants; the validity of the
extrapolation to the smaller sizes is probably the more questionable.
The estimates of labor, utilities, and raw material requirements
(Table E-l) were taken from Reference El. The cost factors were taken
from Appendix B, with two exceptions. The cost of cooling water was
taken to be a capitalized cost for cooling tower water, since the capi-
tal cost of the reduction plant did not include that for a cooling tower.
The credit for low pressure steam was taken to be the same as used by
Allied.
The yield of sulfur in the reduction plant was taken as 92.63 per-
cent, as estimated by Allied.
The estimated annual costs for reduction plants are calculated in
Table E-2 and are presented graphically in Fig. E-2.
E-l
-------
REFERENCE
El. Yodis, A. W. , Applicability of Reduction to Sulfur Techniques to
the Development of New Processes for Removing SO from Flue Gases
— Phase I. Allied Chemical Corporation, Interim Report to
National Air Pollution Control Administration for Period June 1,
1968 - July 31, 1969 (Draft Copy, September 26, 1969), Contract
No. PH 22-68-24
Table E-l
LABOR, UTILITIES, AND RAW MATERIAL REQUIREMENTS
OF SULFUR DIOXIDE REDUCTION PLANTS
Item
Methane
Electricity
Cooling Water
Boiler Feed Water Makeup
Fuel Gas
Catalyst
(Low Pressure Steam)
Labor
Supervision
Quantity
per 1000 Ib S0a input
2940 CF
8.90 kwh
420 gal
5 gal
170 CF
0.18 Ib
(370 Ib)1
54 man-hr/day
8 man-hr/day2
Unit Cost
40C/1000 CF
1.0£/kwh
4
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Table E-2
CAPITAL AND ANNUAL COSTS OF SULFUR DIOXIDE REDUCTION PLANTS
Item
Capital Investment
Annual Cost
A. Depreciation
B. Direct Operating Cost
1 . Labor
2 . Supervision
3. Payroll benefits
4. Maintenance materials
5. Factory supplies
6. Electricity
7 . Methane
8. Fuel gas
9. Cooling water
10. Boiler feed water
11. Catalyst
C. Indirect Costs
1. Controllable
2. Noncontrollable
Gross Total Annual Cost
Steam Credit
Net Total Annual Cost
Cost
Basis
$
$/yr
$/yr
$/yr
$/yr
Sulfur Dioxide Input -IbAr
3,000
342 , 000
22,800
66,830
12,540
19,840
10,260
1,710
2,110
27,940
1,620
400
50
560
44,820
10,260
221,740
2,200
219,540
10,000
695,000
46,330
66,830
12,540
19,840
20,850
3,480
7,050
93,140
5,390
1,330
160
1,850
50,110
20,850
349,750
7,330
342,420
30,000
1,320,000
88 , 000
66,830
12,540
19,840
39,600
6,600
21,140
279,420
16,170
3,990
470
5,560
59,490
39,600
659,250
21,980
637,270
200,000
4 , 000 , 000
266,700
66,830
12,540
19,840
120,000
20,000
140,960
1,862,800
107,800
26,600
3,160
37,060 '
99,690
120,000
2,903,980
146,520
2,757,460
E-3
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10
i
o
I
1
0
0.1
-I I I I I I I
I I I I I I 1 I I
i i i i I 11
I i i i i I I i I
3.000
10.000 100.000 200.000
SULFUR DIOXIDE INPUT — Ih/hr TA-7923-23
FIGURE E-1 CAPITAL COSTS OF SULFUR DIOXIDE REDUCTION PLANTS
E-4
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10
o
TJ
»•*
O
w
|
1
o
z
z
o
1.0
0.1
3.0OO
10,000
SULFUR DIOXIDE INPUT — Ib/hr
100,000 200,000
TA-7923-24
FIGURE E-2 ANNUAL COSTS OF SULFUR DIOXIDE REDUCTION PLANTS
E-5
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