Final Report
Contract No. CPA 22-69-78
FEASIBILITY STUDY OF NEW SULFUR OXIDE
CONTROL PROCESSES FOR APPLICATION
TO SMELTERS AND POWER PLANTS
Part IV: The Wellman-Lord S02 Recovery
Process for Application to Power
Plant Flue 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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1
Final Report
Contract No. CPA 22-69-78
FEASIBILITY STUDY OF NEW SULFUR OXIDE
CONTROL PROCESSES FOR APPLICATION
TO SMELTERS AND POWER PLANTS
Part IV: The Wellman-Lord S02 Recovery
Process for Application to Power
Plant Flue Gases
By: KONRAD T. SEMRAU
Prepared for:
U.S. DEPARTMENT OF HEALTH, EDUCATION, AND WELFARE
NATIONAL AIR POLLUTION CONTROL ADMINISTRATION
SRI Project PMU-7923
Approved:
N. K. HIESTER, Director
Physical Sciences (Materials)
C. J. COOK, Executive Director
Physical Sciences Division
r^apj. STANFORD RESEARCH INSTITUTE
Menlo Park, California 91025 • U.S.A.
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CONTENTS
FOREWORD vii
I INTRODUCTION 1
II OBJECTIVES 5
III SUMMARY 7
IV PROCEDURES 15
A. Formulation of Models 15
B. Cost Factors 16
C. Preparation of Technical Data and Cost Estimates . . 16
V PROCESS DESCRIPTION 17
VI PROCESS DATA AND COST ESTIMATES 27
A. Material Flows 27
B. Capital Cost Estimates 28
C. Operating Cost Estimates 29
VII GENERAL DISCUSSION 41
A. Evaluation of the Wellman-Lord S0o Recovery System . 41
B. By-Product Values 43
C. Variation of Bases of Cost Estimates 45
1. Load Factor 45
2. Amortization Period 46
3. Fixed Charges 47
REFERENCES 49
iii
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APPENDIXES
A. MODELS FOR HYPOTHETICAL POWER PLANTS A-l
B. COST FACTORS USED IN MODEL STUDIES B-l
C. STACKS FOR USE ON CONTROLLED SULFUR DIOXIDE
EMISSION SOURCES C-l
D. ESTIMATION OF THE VALUES OF SULFUR BY-PRODUCTS .... D-l
ILLUSTRATIONS
Figure I Wellman-Lord SO Recovery System . , 20
TABLES
Table I Power Plant Model A — Summary of Estimated Costs
for Wellman-Lord SO2 Recovery System with Contact
Sulfuric Acid Plant 11
Table II Power Plant Model A — Summary of Estimated Costs
for Wellman-Lord SOg Recovery System with Sulfur
Dioxide Reduction Plant 12
Table III Power Plant Model B — Summary of Estimated Costs
for Wellman-Lord SOg Recovery System with Contact
Sulfuric Acid Plant 13
Table IV Power Plant Model B — Summary of Estimated Costs
for Wellman-Lord S02 Recovery System with Sulfur
Dioxide Reduction Plant . 14
Table V Flue Gas Flows in Wellman-Lord System Absorption
Sections 28
Table VI Power Plant Model A — Material Balances for SO2
Recovery System 30
Table VII Power Plant Model B — Material Balances for SO2
Recovery System 32
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TABLES (Concluded)
Table VIII Capital Investments for Wellman-Lord S02 Recovery-
Plants ... 34
Table IX Power Plant Model A — Summary of Capital and
Operating Costs for Wellman-Lord System 35
Table X Power Plant Model A — Summary of Capital and
Operating Costs for Contact Sulfuric Acid Plant. . . 36
Table XI Power Plant Model A — Summary of Capital and
Operating Costs for Sulfur Dioxide Reduction Plant . 37
Table XII Power Plant Model B — Summary of Capital and
Operating Costs for Wellman-Lord System 38
Table XIII Power Plant Model B — Summary of Capital and
Operating Costs for Contact Sulfuric Acid Plant. . . 39
Table XIV Power Plant Model B — Summary of Capital and
Operating Costs for Sulfur Dioxide Reduction Plant . 40
Table D-l Estimated Sulfuric Acid Demand in Selected
Producing Areas, 1966 D-6
Table D-2 Phosphoric Acid Production Costs D-7
v
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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 Plant
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 re-
lease of the reports by Monsanto and Wellman-Lord does not imply that
these companies necessarily concur in all or any of the opinions, judg-
ments, 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
8c 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 evalu-
ation of the technical and economic feasibility of these control processes,
a number of models of smelters were created. Stanford Research Institute
carried out the studies necessary to determine the availability of mar-
kets 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
usefulness 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 applica-
tion 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 S0_ Recovery process. However, both processes are proprietary, and,
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.
1
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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 evaluation of these por-
tions has had to be presented in the form of conclusions without support-
ing 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
verification 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 reasonable.
In other instances, information or estimates were provided by the com-
panies that could not be verified independently or judged for reasonable-
ness; in such cases, the companies have been specifically cited as the
2
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sources. In still other instances, the author did not accept the in-
formation or estimates provided and in some cases substituted his own;
such cases have also been specifically noted.
The cooperating companies, at their own option and through substan-
tial 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 reasonableness
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 re-
view of the results, an effort was made to ensure that the cost estimates
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 themselves.
For convenience, and at the request of DPCE-NAPCA, this final re-
port is presented in four separate and independent parts. This part
deals only with the Wellman-Lord SO Recovery process as applied to
Oi
control of sulfur oxides in the flue gases from power plants.
3
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XI OBJECTIVES
The objectives of the part of the study covered by this report are
as follows:
1. To prepare block flow diagrams of the Wellman-Lord system
showing its configuration and its relation to the boiler
plant 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. From the
estimate of total annual cost, secondary estimates are to be
made of the corresponding incremental cost of producing elec-
tricity, both on the gross basis (without allowance for
by-product recovery credits) and on the net basis (with
allowance for by-product recovery credits).
4. To make a qualitative appraisal of technical constraints on
the application and operation of the control system.
5. To appraise (quantitatively, to the extent permitted by
available data) the economic constraints on the application
of the control system.
6. 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
power plants.
The accomplishment of the objectives is subject to any restrictions
that may be imposed under the terms of the confidentiality agreement be-
tween the Government, Wellman-Lord, Inc. and Stanford Research Institute,
5
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Ill SUMMARY
The Wellman-Lord SO Recovery process is a cyclic absorption-desorp-
dt
tion process lor producing concentrated sulfur dioxide from waste gas
streams, such as power plant flue gas and smelter offgases. It is adapt-
able 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 re-
moved 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 absorbent 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, releas-
ing sulfur dioxide and regenerating the sulfite.
The absorption step is conventional and has been employed commer-
cially as well as studied experimentally. The regeneration step appar-
ently represents 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 stripping 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.
It is essential to the operation of the Wellman-Lord system that
the flue gas entering the absorber be as nearly free of fly ash as
practical. Fly ash collected in the absorbent solution may produce a
7
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variety of mechanical and chemical problems. The electrostatic fly ash
precipitator removes the hulk of the fly ash, but a prescrubber is also
installed immediately upstream from the absorber. The prescrubber col-
lects most of the residual fly ash that escapes collection by the elec-
trostatic precipitator, and also cools and humidifies the flue gas.
The Wellman-Lord system treated in this study is essentially a
conceptual design. Components of the system have been tested indivi-
dually, but a complete, integrated system has not yet been operated.
Another version of the Wellman-Lord process, using potassium sulfite
solution as the absorbent, was tested in a pilot plant and in a demon-
stration plant but was generally unsuccessful. The regeneration system
was basically different from that used with the sodium-base absorbent,
and it did not attain the steam economy originally anticipated. Many
of the problems encountered resulted from entry of fly ash into the
system; the prescrubber used was inadequate in efficiency and plugged
rapidly with fly ash.
The Wellman-Lord system as designed for use with the sodium-base
absorbent appears to be technically feasible and to avoid most of the
problems that were encountered in the demonstration plant using the
potassium-base absorbent system.
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 will 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.
8
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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.
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 absorption tower of the Wellman-Lord system is designed to re-
move 90 percent of the sulfur dioxide from the flue gas. However, some
of the recovered sulfur dioxide is subsequently lost to the atmosphere
from the sulfuric acid plant or the reduction plant used to convert the
sulfur dioxide. The overall sulfur dioxide emission control efficiency
is estimated to be 88 percent for the combination of the Wellman-Lord
system and acid plant, and 84 percent for the Wellman-Lord system and
reduction plant. The sulfur dioxide emissions from the assumed auxiliary
conversion plants were estimated by the author. Various potential methods
might be used to recover part of these emissions. For example, Wellman-
Lord absorbers might be applied to the tail gases, with the spent absor-
bent being returned to the central chemical recovery section that serves
the main absorption section. The acid plant might employ the double-
contact process to increase the sulfur dioxide conversion.
To permit estimates of capital and annual costs of the Cat-Ox and
Wellman-Lord SOg Recovery systems, two models of hypothetical power plants
were created. (Details are presented in Appendix A.) The first (Model A)
is an existing 500-megawatt plant located in central Pennsylvania (Altoona).
The second (Model B) is a new 1000-megawatt plant located on a navigable
river in the Midwest (Cairo, Illinois) . Estimates of annual cost were
based on a plant life of 20 years and on an annual operating time of 7000
hours (load factor 80 percent). The estimated prices for by-product
sulfuric acid are based on an assumed Gulf Coast price for sulfur of
$30/long ton, which is probably as high as can be anticipated in the period
up to 1975.
9
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Cost estimates were made for Wellman-Lord SO^ Recovery systems to
be applied to each hypothetical power plant, and to each recovery plant
were added alternative auxiliary contact sulfuric acid plants and sulfur
dioxide reduction plants. The costs for the auxiliary plants were esti-
mated by the author. The estimates of capital and operating costs for
the four control system models are summarized in Tables I through IV.
The annual costs and the equivalent incremental costs of electricity
generated and fuel used are presented both before and after allowance
for sulfur by-product credits. In the case of the Model A control sys-
tem producing acid, two acid prices are presented in estimating the by-
product credit; the lower of the two is applicable in the event that it
should be necessary to displace existing captive acid producers in the
Pittsburgh area, where the acid from the Altoona plant would have to be
sold.
Although the cost of producing elemental sulfur exceeds that of
producing sulfuric acid, disposing of the acid would often be difficult
in actual cases, whereas the sulfur can be shipped greater distances or
be readily stored.
10
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Table I
POWER PLANT MODEL A
SUMMARY OF ESTIMATED COSTS FOR WELLMAN-LORD SOa RECOVERY SYSTEM
WITH CONTACT SULFURIC ACID PLANT
Item
Cost Basis
Cost
or Credit
Capital Investment
A, Wellman-Lord System
$
7,240,000
B. Contact Acid Plant
$
790,000
Total
$
8,030,000
$ Aw
16.06
Gross Costs Before Credits
$/yr
2,974,165
Mills/kwh
0.850
^/million Btu
8. 50
$/ton of coal
2.26
Credits for Sulfuric Acid
95,025 tons/yr (100$ basis)
$12.00/ton
$ 7.00/ton
1,140,300
665,175
Net Costs After Credits
A. Acid Price = $12/ton
$/yr
1,833,865
Mills/kwh
0.524
^/million Btu
5.24
$/ton of coal
1.40
B. Acid Price = $7/ton
$/yr
2,308,990
Mills/kwh
0.660
^/million Btu
6.60
$/ton of coal
1.76
11
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Table II
POWER PLANT MODEL A
SUMMARY OF ESTIMATED COSTS FOR WELLMAN-LORD SOg RECOVERY SYSTEM
WITH SULFUR DIOXIDE REDUCTION PLANT
Item
Cost Basis
Cost
or Credit
Capital Investment
A. Wellman-Lord System
B. SOjg Reduction Plant
Total
$
$
$
$ Aw
7,240,000
1,040,000
8,280,000
16.56
Gross Costs Before Credits
$/yr
Mills/kwh
^/million Btu
$/ton of coal
3,174,135
0.907
9.07
2.42
Credits for Elemental Sulfur
29,330 short tons/yr
$33.00/long ton
864,190
Net Costs After Credits
$/yr
Mills/kwh
#/million Btu
$/ton of coal
2,309,945
0.660
6.60
1.76
22
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Table III
POWER PLANT MODEL B
SUMMARY OF ESTIMATED COSTS FOR WELLMAN-LORD SOg RECOVERY SYSTEM
WITH CONTACT SULFURIC ACID PLANT
Item
Cost Basis
Cost
or Credit
Capital Investment
A. Wellman-Lord System
B. Contact Acid Plant
Total
$
$
$
$Aw
12,010,000
1,200,000
13,210,000
13.21
Gross Costs Before Credits
$/yr
Mills/kwh
//million Btu
$/ton of coal
4,589,750
0. 656
6.83
1.82
Credits for Sulfuric Acid
182,350 tons/yr (100$ basis)
$10.00/ton
1,823,500
Net Costs After Credits
$/yr
Mills/kwh
//million Btu
$/ton of coal
2,766,250
0.395
4.12
1.10
13
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Table IV
POWER PLANT MODEL B
SUMMARY OF ESTIMATED COSTS FOR WELLMAN-LORD SOa RECOVERY SYSTEM
WITH SULFUR DIOXIDE REDUCTION PLANT
Item
Cost Basis
Cost
or Credit
Capital Investment
A. Wellman-Lord System
B, SOa Reduction Plant
Total
$
$
$
$/kw
12 ,010,000
1,520,000
13 ,530,000
13.53
Gross Costs Before Credits
$/yr
Mills/kwh
^/million Btu
$/ton of coal
4,886,020
0.698
7.27
1.94
Credits for Elemental Sulfur
56,280 short tons/yr
$33.00/long ton
1,658,250
Net Costs After Credits
$/yr
MillsAwh
^/million Btu
$/ton of coal
3,227,770
0.461
4.80
1.28
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IV PROCEDURES
A . Formulation of Models
The formulation of the models of the hypothetical power plants was
based first on conditions specified by DPCE-NAPCA. Part of the remain-
ing model conditions were derived from the NAPCA specifications so as
to be consistent with the latter. Other conditions had to be selected
on a relatively arbitrary basis, and where this was necessary, guidance
was obtained from the literature or from discussions with informed indi-
viduals. In other cases, conditions had to be selected on completely
arbitrary bases, and the author used his own experience and best judgment.
The complete models are presented in Appendix A.
The basic premise employed throughout was that all estimates of the
costs of sulfur oxides control should be made in relation to a base level
represented by the equivalent conventional plant without sulfur oxides
emission controls. Emission control was to be charged with whatever
costs were incurred above those of the conventional, or base-level, plant.
Similarly, it was to be credited with any savings from the cost of the
base-level plant, or with any income derived from sale of by-products.
If the sulfur oxide control system included components common to
the base-level plant, it was to be charged only with the increment in
the costs of these components over those of the corresponding items in
the base-level plant (or credited with the difference if the items in
the conventional plant were more expensive).
The sulfur oxide control system was to be charged with the cost of
power required to move the flue gas through any parts of the system
specific to emission control. It was also to be charged with a propor-
tionate share of the capital cost of the fan and motor if the latter
also supplied draft for the rest of the boiler system.
15
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The model formulations were circulated to DPCE-NAPCA and to Monsanto
Company and Wellman-Lord, Inc. for review before adoption,
B. Cost Factors
The factors used in making cost estimates, and the bases on which
they were adopted, are presented in Appendix B. Salaries, payroll bene-
fits, and overhead were taken to be the same as those used in the previous
15
study of the nonferrous smelting industry.
Estimates of the prices that might be obtained for sulfur by-products
were made by Stanford Research Institute, and are presented in Appendix D.
The choices of the specific sites for the hypothetical plants (within the
general areas specified by DPCE-NAPCA) were made to facilitate estimation
of definite prices.
C. Preparation of Technical Data and Cost Estimates
Wellman-Lord, Inc. prepared the technical designs for the model
control systems, based on the conditions formulated by SRI, and estimated
the capital investments and the utilities and maintenance requirements.
The author reviewed these estimates and accepted most of them. The cost
breakdowns for individual components of the system were supplied by Well-
man-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 sur-
veyed these estimates for general reasonableness and consistency with
the model specifications, but did not attempt further verification.
Some of the original estimates of system requirements or costs
made 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 auxiliary conversion plants (contact sul-
furic acid and sulfur dioxide reduction) were made by the author from
6 28
published data. ' (See also Appendixes D and E of Part II of this repori
16
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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 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. Its product, concentrated sulfur
dioxide, has very limited markets, but can be used in sulfuric acid
manufacture or be reduced to elemental sulfur. 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 gas directly in a con-
Q
tact sulfuric acid plant.
The absorption step in the Wellman-Lord process cycle is conven-
tional, and in closely related forms has been used commercially and been
studied widely on an experimental basis by numerous investigators. Well-
man-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 novel, 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= + S0g + H20 = 2 HSOg"
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In the desorption cycle, the reaction is reversed by application of heat
releasing sulfur dioxide and regenerating the sulfite:
2 HSO ~ = SO = + H O + SO
J u A &
The absorption step has been employed commercially in the pulp and
paper industry (usually with sodium and ammonia as bases) for production
1 5 18
of cooking liquor; ' ' in this case no regeneration step is employed.
The complete cycle, including the desorption step, has been studied ex-
12,13
perimentally and employed occasionally on a limited commercial
12
scale. However, the major economic limitation has been the quantity
of steam required for the absorbent regeneration. In the conventional
13
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
14
absorbent. Johnstone, 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
3 16 IT
are described in Wellman-Lord patents, ' ' One, which uses the
3
potassium-base salts, is described broadly in a Belgian patent, and
27
more specifically in a paper by Watt. It was studied on a pilot plant
26 27
scale and was later investigated in a demonstration plant. The an-
ticipated steam economy 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 regen-
eration process that constitutes the apparently novel feature of the
18
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Wellman-Lord system. Applications have been made for patents, but patent
protection has not yet been received. Consequently, it is not permissible
at this time to describe the basic regeneration process or the system for
carrying it out. The corresponding portions of the regeneration system
are therefore 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 demonstration
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 absorbent
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 re-
action with sulfur dioxide (the actual crystalline compound is the
pyrosulfite instead of the bisulfite), the solution becomes less nearly
saturated. Hence, a portion of the water in the solution can be evaporated
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
absorber and absorbent outlet piping by pyrosulfite crystals may occur.
-------
FIGURE 1 WELLMAN-LQRD S02 RECOVERY SYSTEM FOR
POWER PLANT FLUE GAS
-------
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-
13
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 re-
gardless of the composition of the gas stream treated.
The steam economy of the system can be improved by staging the re-
generation process, but the economic limits of this approach are set by
the concomitant increase in capital investment. The optimum balance be-
tween 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 sul-
fate, 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 esta-
blished; it may, in fact, vary widely depending upon such factors as the
oxygen content of the gas stream, the temperature, and the presence of
contaminants that may act as oxidation catalysts. Wellman-Lord adds
hydroquinone to the absorbent solution as an oxidation inhibitor, but
the effectiveness of this measure is questionable. This point is dis-
cussed 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, the
prescrubber, and the fan. Its size, and hence its cost, is dependent
almost solely upon the volume of gas handled, with only a minor cost
21
-------
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 absorbed 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 oc-
cluded 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. In addition, the particulate
matter may include elements that will act as catalysts for the oxidation
of the sulfite.
Acidic gases and vapors other than sulfur dioxide are also objection-
able, since they will react with and consume absorbent. Some of the sul-
fur trioxide present in the flue gas will be collected in the absorber
and form sodium sulfate, which must then be purged from the system along
with the sulfate formed by the oxidation of part of the sulfite, Hydrogen
chloride and hydrogen fluoride, which may also be present in small quan-
tities, will form sodium chloride and sodium fluoride that must also
eventually be purged from the absorbent, A small quantity of sodium
nitrate may also be formed by the absorption of nitrogen oxides.
In the present study, the model power plants are assumed to have
the same electrostatic fly ash precipitators as would the corresponding
plants without sulfur oxide emission controls. In addition, the pro-
posed Wellman-Lord system includes the prescrubber specified by Wellman-
Lord. The prescrubber, which is only one of several potential alternative
22
-------
types, is provided with trays lor countercurrent gas-liquid contacting.
It is actually constructed within the same shell as the absorber, al-
though this arrangement is not indicated in Fig. 1 and is not essential,
but only a matter of convenience. The gas pressure drop is 4 inches of
water. Wellman-Lord anticipated that the prescrubber will remove up to
about 90 percent of the fly ash present in the flue gas as it comes from
the electrostatic precipitator. On the basis of scrubber performance on
19 20
fly ash reported in the literature, ' the anticipated efficiency ap-
pears reasonable. The water used for scrubbing is recirculated with
makeup water provided to compensate for evaporation and the water in the
fly ash slurry discharged to waste. Since the scrubbing water is recir-
culated, the scrubbed gas should approach saturation at its adiabatic
saturation temperature, which is about 128°F. The cooling and humidifi-
cation of the flue gas will also control the evaporation of water from
the absorbent solution in the absorber.
In addition to collecting most of the fly ash that escapes the
electrostatic precipitator, the prescrubber will also collect a portion
of the sulfur trioxide. However, the sulfur trioxide will react with
water vapor in the prescrubber to form a sulfuric acid aerosol, and the
collection efficiency is not readily predictable; in this study the
author assumed an efficiency of 50 percent. The prescrubber may remove
a portion of any hydrogen chloride and hydrogen fluoride present in the
flue gas, but the acidic, recirculated scrubber water will not be a
favorable absorbent for these gases.
The absorber will also remove additional amounts of the fly ash
and sulfuric acid mist. The author assumed an overall efficiency on
fly ash (for the prescrubber and absorber) of 99.5 percent, which ap-
pears reasonable in view of the total pressure drop of 16 inches of
water resulting from gas-liquid contacting. He also assumed an overall
efficiency of 90 percent on the sulfuric acid mist, but this is much
more speculative.
The flue gas leaving the prescrubber is forced up through the coun-
tercurrent absorbing tower where it contacts the absorbent solution on
23
-------
a series of trays and the sulfur dioxide is absorbed. The treated gas
leaving the top tray passes through an entrainment separator where en-
trained droplets of the absorbent solution are removed, and it is then
discharged to the stack. Because the concentration of the sulfur dioxide
is low, the liquid throughput is also relatively low and special provi-
sions (not indicated in Fig. 1) are made to ensure good gas-liquid contact.
The gas pressure drop across the absorber is about 16 inches of water,
including the entrance and exit losses of both the prescrubber and ab-
sorber. The largest individual absorption towers are rated at 505,000
ACFM of gas, and hence both of the model power plants require multiple
parallel towers (three in the Model A plant and six in the Model B plant).
In the Model B plant the fans and large pumps are driven by back-
pressure steam turbines, whose exhaust supplies the low pressure steam
needed in the chemical recovery section. For the Model A plant, it was
specified in the model conditions (Appendix A) that the fans should be
electrically driven, so that the electrical load is greater than that
for the larger Model B plant.
The rich absorbent solution leaving the bottom of the absorbing
tower is pumped to the chemical recovery section, where the sulfur diox-
ide 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 to the main condenser, which is operated at about 200°F. Most of
the steam is condensed, but the temperature is high enough that little
of the sulfur dioxide is absorbed in the condensate, which can be re-
turned to the system as makeup water. The hot, wet sulfur dioxide stream
leaving the receiving tank of the main condenser enters the second con-
denser, where it is cooled to about 120°F. Most of the remaining water
vapor condenses, but the condensate contains a high concentration of
dissolved sulfur dioxide. This second condensate stream must therefore
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 120°F and
saturated with water vapor, which then constitutes less than 12 percent
24
-------
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 pre-
vent 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.
The fly ash that is collected in the absorber must be purged from
the system. Hence, a sidestream of the lean absorbent solution return-
ing to the absorber is diverted through a centrifuge where the fly ash
is removed and the clarified solution is returned to the absorber circuit.
The amount of fly ash to be removed at this point and, hence, the size
of the sidestream and centrifuge, is obviously dependent upon the effi-
ciency of the gas cleaning system ahead of the absorber.
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 sul-
fur 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 carbon-
ated to precipitate 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 recovery system might be about $12 to $15 per ton of sodium hydrox-
ide recovered, including the cost of lime, makeup sodium hydroxide, opera-
tion, and depreciation of equipment. The higher figure, $15/ton, is
assumed in this study.
25
-------
In the purging of sulfate, a portion of the sulfur dioxide is lost
because of purging of the associated sulfite. The author calculated the
sulfur dioxide loss as 9.5 percent of that absorbed from the flue gas,
assuming that 1.0 percent is oxidized and that 40 percent of the sulfur
trioxide in the flue 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.
26
-------
VI PROCESS DATA AND COST ESTIMATES
A. Material Flows
The estimated quantities of coal burned in the two hypothetical
power plants — and hence the calculated quantities of flue gas, sulfur
compounds, and fly ash to be handled — were fixed by the assumed heat
rates (see Appendix A). The heat rates assumed for the Model A and
Model B power plants were 10,000 and 9,600 Btu/kwh, respectively. These
values are probably between 5 and 10 percent higher than those that might
reasonably be obtained at comparable actual plants. Although the assumed
heat rates are perhaps excessively conservative, they are well within the
probable precision of the other estimates made for the study.
Although the calculated volume of flue gas generated was increased
by use of the high values of heat rate, no allowance was made for any
increase in gas volume resulting from air infiltration into the flue gas
handling system, which occurs in actual practice.
The estimated rates and conditions of gas flow through the Model A
and Model B Wellman-Lord absorbing sections are presented in Table V.
The treated flue gas leaving the absorber will have a water vapor content
somewhat below the saturation value because of the vapor pressure lower-
ing by the salts in the absorbent solution. The author assumed that
there is no change in the moisture content of the gas after it leaves
the prescrubber, saturated at 128°F.
The material balances for sulfur, fly ash, and absorbent chemicals
are presented in Tables VI and VII, which include the sulfur balances
for the alternative sulfur dioxide conversion plants (contact sulfuric
acid and sulfur dioxide reduction). The performance of the prescrubber
and absorber on sulfur trioxide and fly ash was based largely on esti-
mates or assumptions by the author. The sulfur trioxide emitted to the
atmosphere would actually be in the form of sulfuric acid mist. The
original estimates of sodium hydroxide entrainment loss provided by
27
-------
Wellman-Lord were tenfold higher than the values adopted by the author,
who believes that the values given in Tables VI and VII should be reason-
able unless reentrainment should take place in the eliminator.
Table V
FLUE GAS FLOWS IN WELLMAN-LORD
SYSTEM ABSORPTION SECTIONS
Gas Stream
Temp
(°F>
Water Vapor
Content
(*)
Gas Flow Rate
(acfm)
Model A
Model B
Entering
Prescrubber
325
8.08
1,564,360
3,003,230
Leaving
Absorber
130
14.4
1,262,710
2,424,140
B . Capital Cost Estimates
The estimates of capital investment for the sulfur dioxide recovery
plants presented in Table VIII, IX, and XII were prepared by Wellman-Lord
with three exceptions. The author added allowances to each plant model
for ductwork connecting the outlets of the absorbers to the stacks. The
original estimates did not clearly indicate inclusion of such estimates.
The author also provided the credit for the stack for the Model B
plant, which represents the difference between the cost of an 800-foot
stack used with no sulfur oxide emission controls, and that of a 500-foot
stack used when emission controls are applied (see Appendix C). The same
stack credit was assigned arbitrarily to both the Cat-Ox and Wellman-Lord
control systems for the Model B plant. The precision of the estimate is
low, and the credit is provided mostly for purposes of illustration. As
is discussed in Appendix C, some reduction of stack height in this model
4
case appears justified. It appears to the author, as to some others,
that the problems of securing dispersion of scrubbed stack gases have
been overstressed.
28
-------
The estimates of capital investment for the contact sulfuric acid
plants (Tables X and XIII) were taken from the curve of Fig, D-l, Appen-
dix D of Part II of this report. The curve was derived from the data of
Connor,6 and gives only the cost of the acid-making section of a contact
plant; since the feed to the plant consists of concentrated sulfur diox-
ide, no gas-producing section is required. The plants were sized for
the total amount of sulfur dioxide absorbed by the Weilman-Lord systems
without allowance for losses.
The estimates of capital investment for the sulfur dioxide reduction
plants (Tables XI and XIV) were based on an estimate made for a concep-
28
tual design produced by Allied Chemical Corporation. The specific
values were taken from the curve of Fig, E-l, Appendix E of Part II of
this report. The plants were sized for the total amount of sulfur diox-
ide absorbed by the Wellman-Lord systems, without allowance for losses.
C• Operating Cost Estimates
The estimates of electrical power, steam, water, and labor were
developed by Wellman-Lord (Tables IX and XIII), and the corresponding
operating costs were calculated by application of the appropriate cost
factors. The ratio of sodium hydroxide and hydroquinone makeup to sodium
sulfate purged was supplied by Wellman-Lord, but the author made his own
estimates of the amount of sulfate formed, as is discussed above in Sec-
tion V,
Wellman-Lord did not provide an estimate of annual maintenance cost,
but the author adopted the figure of 3 percent of capital investment used
generally for other systems in this study.
The operating requirements for the contact sulfuric acid plants were
6 8
adopted from ranges of values presented in the litrrature, ' Those for
the sulfur dioxide reduction plants were taken fro.n the report by Allied
28
Chemical Corporation.
29
-------
Table VI
POWER PLANT MODEL A
MATERIAL BALANCES FOR SQg RECOVERY SYSTEM
Item
Quantity
(lb/hr)
Percent
of
Input
I. Wellman-Lord System
A. Sulfur Dioxide
1. Input
22,200
100
2. Absorbed
19,980
90
3. Emitted
2,220
10
4. Oxidation and purge loss
1,893
8.52
5. Recovered
18,090
81.48
B. Sulfur Trioxide
1. Input
391.0
100
2. Collected in prescrubber
195,5
50
3. Collected in absorber
156.4
40
4. Emitted
39.1
10
C. Fly Ash
1. Input
1,405
100
2. Collected in prescrubber
1,265
90
3. Collected in absorber
134
9.5
4. Emitted
7.0
0.5
D. Sodium Hydroxide Losses
1, Entrainment
12.6
2, Reaction with S03
971
3. Oxidation of SOa
1,552
E. Hydroquinone Loss
4.6
(continued)
30
-------
Table VI (Concluded)
Item
Quanti ty
(lb/hr)
Percent
of
Input
II. Sulfuric Acid Plant
A. Sulfur Dioxide
1. Input
2. Converted to acid
3. Emitted
B. Sulfuric Acid Produced
18,090
17,730
362
27,150
100
98
2
III. Sulfur Dioxide Reduction Plant
A. Sulfur Dioxide
1. Input
2. Converted to sulfur
3. Emitted
B. Sulfur Produced
18,090
16,760
1,330
8,380
100
92.63
7,37
IV. W-L System plus Acid Plant
A. Sulfur Dioxide
1. Emitted
2. Usefully recovered
3. Lost in purge
2,582
17,730
1,893
11.63
79.85
8.52
V. W-L System plus Reduction Plant
A. Sulfur Dioxide
1. Emitted
2. Usefully recovered
3. Lost in purge
3,550
16,760
1,893
16.00
75.48
8.52
31
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Table VII
POWER PLANT MODEL B
MATERIAL BALANCES FOR S03 RECOVERY SYSTEM
Item
Quantity
(lb/hr)
Pe rcen t
of Input
I. Wellman-Lord System
A. Sulfur Dioxide
1. Input
42,619
100
2. Absorbed
38,360
90
3, Emitted
4,262
10
4. Oxidation and Purge Loss
3,634
8. 52
5. Recovered
34,720
81.48
B. Sulfur Trioxide
1. Input
750,7
100
2. Collected in prescrubber
375. 3
50
3. Collected in absorber
300.3
40
4. Emitted
75.1
10
C. Fly Ash
1 . Input
269.8
100
2. Collected in prescrubber
242.8
90
3, Collected in absorber
25.6
9.5
4. Emitted
1.35
0. 5
D. Sodium Hydroxide Losses
1. Entrainment
24,2
2. Reaction with S03
1,864
3. Oxidation of S03
2,980
E. Hydroquinone Loss
8.8
(continued)
32
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Table VII (concluded)
Item
Quantity
( lb/hr)
Percent
of Input
II.
Sulfuric Acid Plant
A.
Sulfur Dioxide
1 . Input
34,720
100
2. Converted to Acid
34,030
98
3, Emitted
694
2
B.
Sulfuric Acid Produced
52,100
III.
Sulfur Dioxide Reduction Plant
A.
Sulfur Dioxide
1. Input
34,720
100
2. Converted to Sulfur
32,160
92.63
3. Emitted
2, 559
7.37
B.
Sulfur Produced
16,080
IV.
W-L
A.
System plus Acid Plant
Sulfur Dioxide
1.. Emitted
4,956
11.63
2. Usefully Recovered
34,030
79.85
3. Lost in Purge
3,634
8. 52
V.
W-L System plus Reduction Plant
A.
Sulfur Dioxide
1. Emitted
6,281
16. 00
2. Usefully Recovered
32,160
75.48
3. Lost in Purge
3,634
8. 52
33
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Table VIII
CAPITAL INVESTMENTS FOR WELLMAN-LORD
SOs RECOVERY PLANTS1
Model A
Absorber System
$ 3 ,640,000
Chemical Recovery System
3,600,000
Total
$ 7,240,000
$14. 50 Aw
Model B
Absorber System
$ 7,460,000
Chemical Recovery System
5,400,000
Total
$12,860,000
Credit for Stack Height Reduction
850,000
Net total
$12,010,000
$12.00/kw
1 Not including auxiliary plants for conversion
of concentrated sulfur dioxide. Capital in-
vestment does not include land, spares, interest
on investment during construction, start-up
expense, working capital, or royalty on process.
34
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Table IX
POWER PLANT MODEL A
SUMMARY OF CAPITAL AND OPERATING COSTS OF WELLMAN-LORD SYSTEM
Item
Quantity
Cost Basis
Cost
Capital Investment
$
7,240,000
Annual Costs
A. Fixed Charges
14.5$ of capital investment
1,049,800
B. Direct Operating Costs
1. Operating labor
25,667 hr/yr
$3.75/hr
96,250
2. Supervi sion
7,000 hr/yr
$4.75/hr
33,250
3. Payroll benefits
25^ of labor + supervision
32,375
4. Maintenance
3% of capital investment
217,200
5. Electricity
8,500 kw
0.55^/kwh
327,400
6. Steam
140,000 lb At
60.2^/1000 lb
589,960
7. Scrubbing water
550 gpm
2^/1000 gal
4,620
8. Process water
14,000 lb/hr
4.8^/1000 lb
4,700
9. Cooling water
5,500 gpm
2^/1000 gal
46,200
10. Sodium hydroxide—makeup
12.6 lb/hr
$65/ton
2,870
11. Sodium hydroxide—recovered
2,523 lb/hr
$15/ton
132,460
12. Hydroquinone
4.6 lb/hr
82?f/lb
26,400
C. Indirect Costs
1. Overhead
50^ of labor + supervision + maintenance
173,350
Total Annual Cost (Gross)
$/yr
2,736,835
Mills/kwh
0.782
^/million Btu
7.82
$/ton of coal
2.08
-------
Table X
POWER PLANT MODEL A
SUMMARY OF CAPITAL AND OPERATING COSTS FOR CONTACT SULFURIC ACID PLANT
Item
Quantity
Cost Basis
Cost
Capital Investment
$
790,000
Annual Costs
A. Fixed Charges
14.5% of capital investment
114,550
B. Direct Operating Costs
1. Operating labor
7,000 hrs/yr
$3.75/hr
26,250
2, Supervision
1,750 hrs/yr
$4.75/hr
8,310
3. Payroll benefits
25% of labor + supervision
8,640
4. Maintenance
3% of capital investment
23,700
5. Electricity
25 kwh/ton acid
0.55£/kwh
13,070
6. Process water
20 gal/ton acid
20£/l,000 gal
380
7. Cooling water
7,000 gal/ton acid
2
-------
Table XI
POWER PLANT MODEL A
SUMMARY OF CAPITAL AND OPERATING COSTS FOR SULFUR DIOXIDE REDUCTION PLANT
Item
Quantity
Cost Basis
Cost
Capital Investment
$
1,040,000
Annual Costs
A. Fixed charges
14.5% of capital investment
150,800
B. Direct Operating Costs
1. Operating labor
54 man-hr/day
$3.75/hr
59,060
2. Supervision
8 man-hr/day
$4.75/hr
11,080
3. Payroll benefits
25% of labor + supervision
17,540
4. Maintenance
3% of capital investment
31,200
5. Electricity
161 kw
0.55£/kwh
6,200
6. Methane
53,200 CF/hr
30
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Table XII
POWER PLANT MODEL B
SUMMARY OF CAPITAL AND OPERATING COSTS FOR TOLLMAN-LORD SYSTEM
I tem
Quantity
Cost Basis
Cost
Capital Investment
$
12,010,000
Annual Costs
A. Fixed Charges
14 .5^ of capital investment
1,741,450
B. Direct Operating Costs
1. Operating labor
25,667 hr/yr
$3.75/hr
96,250
2. Supervision
7,000 hr/yr
§4.75/hr
33,250
3. Payroll benefits
25^ of labor + supervision
32,375
4. Maintenance
3^ of capital investment
360,300
5. Electricity
4,960 kw
0.55^/kwh
191,000
6. Steam
270,000 lb/hr
60.2^/1000 lb
1,137,780
7. Scrubbing water
1,100 gpm
2^/1000 gal
9,240
8. Process water
27,000 lb/hr
4.8(271000 lb
9,072
9. Cooling water
10,000 gpm
2^/1000 gal
84,000
10. Sodium hydroxide—makeup
24.2 lb/hr
$65/ton
5,505
11. Sodium hydroxide—recovered
4,844 lb/hr
$15/ton
254,310
12. Hydroquinone
8.8 lb/hr
82^/lb
50,510
C. Indirect Costs
1. Overhead
504 of labor + supervision + maintenance
244,900
Total Annual Cost (Gross)
$/yr
4,249,942
Mills/kw'h
0.607
(^/million Btu
6.32
!r/ton of coal
1.69
-------
Table XIII
POWER PLANT MODEL B
SUMMARY OF CAPITAL AND OPERATING COSTS FOR CONTACT SULFURIC ACID PLANT
Item
Quantity
Cost Basis
Cost
Capital Investment
$
1,200,000
Annual Costs
A. Fixed Charges
14.5% of capital investment
174,000
B. Direct Operating Costs
1. Operating labor
7,000 hr/yr
$3.75/hr
26,250
2. Supervision
1,750 hr/yr
$4.75/hr
8,310
3. Payroll benefits
25% of labor + supervision
8,640
4. Maintenance
3% of capital investment
36,000
5. Electricity
25 kwh/ton acid
0.55£/kwh
25,070
6. Process water
20 gal/ton acid
20^/1,000 gal
730
7. Cooling water
7,000 gal/ton acid
2
-------
Table XIV
POWER PLANT MODEL B
SUMMARY OF CAPITAL AND OPERATING COSTS FOR SULFUR DIOXIDE REDUCTION PLANT
Item
Quantity
Cost Basis
Cost
Capital Investment
$
1,520,000
Annual Costs
A. Fixed Charges
14.5% of capital investment
220,400
B. Direct Operating Costs
1, Operating labor
54 man-hr/day
$3.75/hr
59,060
2. Supervision
8 man-hr/day
$4.75/hr
11,080
3. Payroll benefits
25% of labor + supervision
17,540
4. Maintenance
3% of capital investment
45,600
5. Electricity
309 kw
0.55£/kwh
11,900
6. Methane
102,100 CF/hr
30^/1,000 CF
214,410
7. Fuel gas
5,900 CF/hr
30£/l,000 CF
12,390
8. Cooling water
14,580 gal/hr
2£/l,000 gal
2 ,040
9. Boiler feed water
174 gal/hr
40
-------
VII GENERAL DISCUSSION
A. Evaluation of the Weliman-Lord S0o Recovery System
The Wellman-Lord process using sodium-base absorbent, described
above, appears to be technically feasible. The regeneration of the ab-
sorbent, which is, in the writer's opinion, the only really novel aspect
of the process, appears to be feasible and to offer a significant reduc-
tion 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 reduc-
tion of twofold or threefold in the steam requirement.
The use of the potassium-base absorbent should also be technically
feasible, but the most recent information available indicates that the
economic advantages originally anticipated by Wellman-Lord are not
realizable.
Serious operating problems were encountered in the demonstration
plant using the potassium-base absorbent. Most of these were of a
mechanical nature and not related to inherent defects in the basic
absorption system. However, they had a critical effect on the testing
program. The most serious problem was evidently the plugging of the
prescrubber (and sometimes parts of the absorber system) by fly ash.
The prescrubber was defective in mechanical design and inadequate in
efficiency. However, the prescrubber proposed by Wellman-Lord for use
in the Model A and B power plant control systems is an established de-
sign that has been widely used, and it should provide higher efficiency
as well as fewer problems with plugging. Even should this prescrubber
design not prove entirely satisfactory from the standpoint of plugging,
there are other designs available that should be superior in this respect
21 22
and should deliver the same efficiency at the same pressure drop. '
41
-------
These alternative scrubber designs should also be comparable in cost.
Be I ore the final design of prescrubber is chosen for a large recovery
plant (commercial or demonstration), operating tests should be made
with small pilot plant units.
The use of the high-efficiency electrostatic precipitators to col-
lect the bulk of the fly ash is probably desirable, if not a practical
necessity, so long as only a low-energy type of prescrubber is employed.
If the precipitator should be omitted entirely or should have an effi-
ciency of no more than 90 percent, the prescrubber should be of the
high-energy, high-efficiency type. The reported performance of scrubbers
on fly ash from boilers firing pulverized coal19'^° indicates that the
1ly ash is actually relatively easy to collect, and that high collection
efficiencies can be attained with moderate power inputs (6 to 10 inches
of water pressure drop).
In the demonstration plant some plugging problems were encountered
at the bottom of the absorption tower because of the formation of potas-
sium pyrosulfite crystals. The problem arose because the pyrosulfite is
the less soluble of the two potassium salts and reached its highest con-
centration at the bottom of the tower. In the sodium-base system, where
the pyrosulfite is the more soluble of the two salts, such a problem is
more easily avoided unless excessive evaporation of water from the solu-
tion by the flue gas should take place.
The type of absorber proposed by Wellmari-Lord is neither the only,
nor necessarily the best, that might be employed. Stone & Webster En-
24
gineering Corp. prepared estimates of the costs of various types of
absorbers for use with the same sodium sulfite-bisulfite system. The
estimated costs of the absorber systems ranged from $2.25 to $4.76/ACFM,
based on exit gas conditions. The estimated costs of the Wellman-Lord
absorber systems for Model A and Model B are $2.88 and $2.08/ACFM, re-
spectively. Except for the factor of cost, however, Stone & Webster
was unable to present any information that provides a clear choice be-
tween absorber designs.
42
-------
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 generally
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 con-
2 5 9 14 18
fidence. Experience with other recovery systems ' ' ' ' suggests that
the amount of the sulfur dioxide oxidized might be in the range of 0.2 to
2.0 percent, but could possibly be as much as 5 to 10 percent or even
higher, dependent upon the conditions of the particular case. Two of the
2, 18
gas £
10,11,23
2 18
most critical factors are the oxygen concentration in the gas ' and the
presence of oxidation catalysts in the absorbent solution.
High efficiency cleaning of the flue gas ahead of the absorber will
be helpful because it will reduce the introduction of potential oxidation
catalysts into the absorbent. However, some of the catalysts are effec-
tive 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. There-
fore, 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.
B. By-Product Values
Over the period of about one year (1969), sulfur supplies have
passed abruptly from shortage to surplus, and it is now difficult if
not impossible to determine whether there is an established price for
sulfur. The Gulf Coast price of sulfur produced by the Frasch process
has long been the base level for world sulfur prices, but it is appar-
ently not so at present, and it is uncertain whether it will be again.
15
In the previous study of the nonferrous smelting industry it was
estimated that the average Gulf Coast price of sulfur over the period
43
-------
to 1975 might be $30/long ton. When the estimate was made, the Gulf
Coast price was over $40/long ton; currently sulfur is reported to be
selling in some areas for less than $15/long ton. Although sulfur
prices could rise again during the next five years, the figure of $30
now appears to be at the optimistic end of the probable range of prices
that may prevail to 1975 (see Appendix D) . Nevertheless, it has been
used as the basis for by-product price estimates in this study because
there appears to be none that can be used with significantly greater
confidence, and the results still have some usefulness if their limi-
tations are understood.
Even the original estimate of $30/long ton f.o.b. the Gulf Coast
was based on the assumption that no very substantial quantity of sulfur
would be recovered at power plants before 1975. The amounts of sulfur
and sulfur by-products potentially recoverable from the coal consumed
by power plants (whether from the coal itself or from the flue gases)
are so large that recovery of a substantial part of them could completely
disrupt existing local sulfur markets. The estimates made of sulfur and
sulfuric acid prices that might be obtained by the Model A and Model B
power plants (see Appendix D) embody the tacit assumption that no other
large power plants near the assumed locations will also be recovering
sulfur by-products.
In summary, it appears that the credits assigned to sulfur by-
products from the Model A and Model B plants are the highest that can
be reasonably expected in the period up to 1975 . The actual credits
might be much lower, and the situation that may exist after 1975 is
unpredictable on the basis of currently available information. Conser-
vatively, the long-term assessment of the economics of recovery processes
should probably be based on the gross costs without allowance for by-
product credits.
It is sometimes suggested that recovery processes can be more eco-
nomically applied to power plant flue gases if coals with extra high
sulfur contents are used. Since the sulfur dioxide concentration in
the flue gas will be increased, the unit cost of production of the sulfur
44
-------
by-product should indeed be reduced. However, if the by-products must
enter markets that are already glutted, the additional quantity of
material may merely exert an increased downward pressure on prices.
The actual advantage, if any, of using the high sulfur coal would
probably be related to lower purchase prices for the coal.
C. Variations of Bases of Cost Estimates
The choice of some of the economic factors assumed in making the
cost estimates was to a considerable degree arbitrary, as it has been
in similar studies made by other workers. There is no general agreement
among different workers on what constitutes a realistic set of bases.
The result is that estimates for different control systems have been
made on a wide variety of bases — frequently only vaguely defined —
so that the results cannot be directly compared. In this study, it has
been assumed that valid comparison of different control systems is more
important than absolute accuracy, and hence that it is more important
that the bases used should be uniform than that they should be the "cor-
rect" ones — whatever the latter might be. An effort has been made to
describe fully all bases used, so that the results can be recalculated
to any different bases that other persons may wish to use. Although
the bases chosen for the models (Appendixes A and B) are believed to
be generally reasonable, it must be noted that other possible assumptions
regarding some factors might be equally reasonable but lead to markedly
different cost estimates.
1. Load Factor
The assumed annual load factor of 80 percent (equivalent to 7,000
hours per year of operation at full load) was specified by DPCE-NAPCA.
It appears to be fairly reasonable for large, new, or relatively new
baseload plants such as those typified by Models A and B. It is less
certain whether the 80 percent factor can be assumed to hold over the
whole of the assumed life (20 years) of the emission control systems.
The Federal Power Commission assumed (Hydroelectric Power Evaluation,
Report P-35, March 1968) that the annual load factor will be relatively
high over the first half of the estimated 30- to 35-year service life
45
-------
of a power plant, then will diminish, giving a lifetime average of 55 to
7
GO percent. Dennis and Bernstein assumed a load factor of 60 percent
over an assumed 11-year life for the emission control systems. Although
their assumptions may be somewhat too conservative, the author believes
that the average factor of 80 percent assumed in the present study is
probably too optimistic.
Because of the high fixed charges associated with the emission
control systems, the load factor exerts great leverage in determining
the incremental cost of producing electricity. The estimates of incre-
mental unit costs presented in Tables I through IV are probably as low
as they can be expected to be with respect to load factor. The actual
costs would probably be higher.
2. Amortization Period
Depreciation for the emission control systems was based on an as-
sumed useful life of 20 years. It was considered that any control
system that might go into service at the present time will be obsolete
within 20 years, if not in a shorter period. It was assumed that tax
laws and the regulations of public utility commissions will permit the
use of this period in determining taxes and rates for electrical power.
At this time there appear to be no precedents to indicate how the exist-
ing depreciation schedules of the tax laws might be interpreted in some
7
specific cases. Dennis and Bernstein assumed a depreciation period of
11 years, as specified for chemical plants by the Internal Revenue Ser-
vice. Their choice of depreciation period apparently was influenced by
a conviction that the recovery processes would become obsolete within
the 11-year period as well as by the IRS regulations.
Some components of sulfur dioxide recovery systems are relatively
conventional items of chemical plant equipment. For example, a basi-
cally standard contact process plant would be used to convert concen-
trated sulfur dioxide from a recovery system to sulfuric acid. Allowances
for depreciation of the acid plant might then depend upon whether the
plant was owned by the power company and considered to be part of the
utility, or owned by an adjacent chemical company and considered to be
part of a chemical plant.
46
-------
For the purposes of the present study, such considerations were
ignored and it was assumed that all components of the control and re-
covery systems would be depreciated over the 20-year period. Neverthe-
less, it must be recognized that the assigned useful life of the control
system, like the load factor of the power plant, can exert a very great
weight in determining the incremental unit cost of producing electricity.
The author believes that, on the whole, the assumption of a useful life
of 20 years is as optimistic as can now be made for any control system,
and that the actual useful lifetimes might well prove to be shorter. In
such case, the incremental unit costs of producing electricity given in
Tables I through IV could be substantially increased.
3 . Fixed Charges
The total fixed charge of 14.5 percent of the capital investment
25
per year was adopted from a report by the Tennessee Valley Authority,
which was in turn adapted from guidelines suggested by the Federal Power
Commission, and embodies the assumption of a 20-year depreciation period.
Recent increases in both taxes and the cost of capital are producing
rises in the fixed charges associated with electric power generation.
The figure of 14.5 percent is therefore lower than may be appropriate
at the present time. It has been used in the present study primarily
for the sake of consistency with previous studies, and because the un-
certainties in other factors assumed appear to make greater elaboration
of the capital charge estimates of limited utility. The trend in taxes
and capital costs will, however, tend to increase the incremental unit
costs of electrical power generation estimated in Tables I through IV.
47
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1
2
3
4
5
6
7
8
9
10
11
12
13
REFERENCES
Aho, William A. , The Jenssen Exhaust Scrubber — An Effective Air
Protection System, Tappi 52^ (4), 620-623 (Apr. 1969)
Applebey, M. P., The Recovery of Sulfur from Smelter Gases, J.
Soc. Chem. Inc. 56, 139T-146T (May 1937)
Beckwell Process Corporation, Process for Recovery of Sulfur Dioxide,
Belgian Pat. No. 706,449 (May 13, 1968)
Boyer, A. E., F. B. Kaylor, T. V. Ward, and F. J. Gottlich,
Atmospheric Dispersion of Saturated Stack Plumes, Des. Oper. Air
Pollut. Contr., Pap, MECAR Symp. 1968 (Pub. 1969), 32-40
Clement, J. L., and W. L. Sage, Ammonia-Base Liquor Burning and
Sulfur Dioxide Recovery, Tappi 52 (8), 1449-1456 (Aug. 1969)
Connor, J,. M. , Economics of Sulfuric Acid Manufacture, Chem. Eng.
Progr. 64 (11), 59-65 (Nov. 1968)
Dennis, R., and R. H. Bernstein, Engineering Study of Removal of
Sulfur Oxides from Stack Gases, Report No. GCA-TR-68-15-G, American
Petroleum Institute, New York, N.Y. (1968)
Duecker, W. W. , and J. K. West (Eds.), "The Manufacture of Sulfuric
Acid," Reinhold Publishing Co., New York (1959)
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)
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)
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)
Johnstone, H. F., Recovery of Sulfur Dioxide from Dilute Gases,
Pulp Paper Mag. Can. 53 (4), 105-112 (Mar. 1952)
Johnstone, H. F., H. J. Read, and H. C. Blankmeyer, Recovery of
Sulfur Dioxide from Waste Gases. Equilibrium Vapor Pressure over
Sulfite-Bisulfite Solutions. Ind. Eng. Chem. 30 (1), 101-109
(Jan. 1938)
49
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14. 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)
15. 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
16. 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)
17. Miller, Leo A., and Jack D. Terrana (to Wellman-Lord, Inc.),
Process for Recovering Sulfur Dioxide from Gases Containing Same,
U.S. Pat. No. 3,485,581 (Dec. 23, 1969)
18. 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)
19. Plumley, A. L. , J. Jonakin, J. R. Martin, and J. G. Singer, Removal
of S02 and Dust from Stack Gases -- A Progress Report on the C-E
Air Pollution Control System, Combustion 40 (1), 16-23 (July 1968)
20. Pollock, W. A., J. P. Tomany, and Garry Frieling, Flue-Gas Scrubber,
Mech. Eng. 89 (8), 21-25 (Aug. 1967)
21. Semrau, K. T. , Correlation of Dust Scrubber Efficiency, J. Air
Poll. Control Assoc. 1_0 (3), 200-207 (June 1960)
22. Semrau, K. T., Dust Scrubber Design -- A Critique on the State of
the Art, J. Air Poll. Control Assoc. 13 (12), 587-594 (Dec. 1963)
23. Srivastava , R. D., A. F. McMillan, and I. J. Harris, The Kinetics
of Oxidation of Sodium Sulfite, Can. J. Chem. Eng. 46 (3), 181-184
(June 1968)
24. Stone & 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
25. Tennessee Valley Authority, Sulfur Oxide Removal from Power Plant
Stack Gas -- Use of Limestone in Wet-Scrubbing Process, Conceptual
Design and Cost Study No. 2 for National Air Pollution Control
Administration (1969)
26. 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)
50
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27. Watt, Stuart G., Wellman-Lord SO^ Recovery Process, Wellman-Lord,
Inc., Lakeland, Florida (1969)
28. 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.
51
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Appendix A
MODELS FOR HYPOTHETICAL POWER PLANTS
For the hypothetical power plant models, DPCE-NAPCA specified the
following conditions:
1. The two model plants shall be an existing 500-MW plant
located in central Pennsylvania and a new 1000-MW plant
located in the Midwest on a large navigable river.
2. Both plants shall have dry-bottom boiler furnaces and
shall use coal containing 3 percent of sulfur and 9 percent
of ash. The flue gases shall contain 0.21 percent of
sulfur oxides.
3. The base load shall be 7,000 hours per year (load factor
80 percent).
4. The existing plant shall have an economizer with an exit
gas temperature of 750°F and an electrostatic precipitator
operating at 325°F. The new plant shall have no temperature
constraints on gas cleaning equipment. The height of the
stack for the existing plant shall be 500 feet. Sufficient
space shall be available to permit installation of any
equipment needed.
5. The efficiency of removal of the sulfur dioxide shall be
90 percent. Regional air and water pollution regulations
shall be applicable.
DPCE-NAPCA also supplied an analysis of the coal assumed to be
burned and a calculated analysis of the flue gas formed:
A-l
-------
Analysis of Coal, As Fired
Constituent lb/100 lb, Coal as Fired
C (burned) 72.7
H2 5-1
N2 1.4
S 3.0
°2 6*9
h2o 1.0
Ash 9.0
100.0
Flue Gas Analysis^
Component Percent by Volume
C0„ 13.89
S0„ and SO 0.2148
m J
O 3.301
N2 74.514
H2° 8.08
100.0
Excess air used = 20 percent. Moisture
content of air = 0.013 lb/lb dry air.
A-2
-------
Quantity of Flue Gas Formed
Basis Lb Moles Gas/100 lb Coal Burned
Wet 43.67
Dry 40.14
The remainder of the model conditions were derived by SRI from the
foregoing conditions, supplemented by assumptions based on data in the
literature, or on the experience or judgment of the author.
A. Conditions Common to Both Model Plants
1. Heating Value of Coal
13,325 Btu/lb (calculated from DuLong Formula)
2. Flue Gas Generated
Volume of gas per 100 lb of coal burned (wet basis)
= 16,880 SCF (70°F, 1 atm)
= 25,013 CF (325°F, 1 atm)
= 43,335 CF (900°F, 1 atm)
3. Air Inleakage into Flue Gas Handling System
Assumed nil
4. Fly Ash
Ash = 9.0 lb/100 lb coal burned
Unburned carbon = 0.35 lb/100 lb coal burned
Total ash and unburned carbon = 9.35 lb/100 lb coal burned
Fly ash (ash + carbon) leaving furnace in flue gas
= 80 percent of total
= 7.49 lb/100 lb coal burned
Fly ash concentration in flue gas = 3.105 grains/SCF (70°F,
1 atm, wet basis)
5. Sulfur Oxides Concentrations
S0_ + SO, = 0.2148 percent by volume, wet basis
« 3
SO = 30 ppm = 0.003 percent by volume, wet basis
o
S0o = 0.2118 percent by volume, wet basis
A-3
-------
6. Sulfur Emission
All of sulfur assumed to go into flue gas
Total sulfur emitted = 3.0 lb/100 lb coal burned
Sulfur emitted as SO^
= 0.0417 lb S/100 lb coal burned
= 0.1042 lb SO /100 lb coal burned
Sulfur emitted as S02
= 2.958 lb S/100 lb coal burned
= 5.916 lb S02/100 lb coal burned
7. Efficiency of Emission Control
Sulfur oxide removal efficiency = 90 percent
Allowable discharge of sulfur (as SO^, SO^, or HgSO^)
not to exceed 10 percent of the input sulfur to the
collection system, or 0.3 lb S/100 lb coal burned
8. Fuel Available for Reheating Flue Gas
No. 2 fuel oil (sulfur content 0.35 percent or lower)
9. Reductant Available for Reducing Sulfur Dioxide
Natural gas (heating value 1000 Btu/CF)
B. Model A Plant
1. Location
Altoona, Pennsylvania
2. Generating Capacity
500 MW
3. Temperature of Flue Gas
At exit from economizer: 750°F
o
At electrostatic precipitator: 325 F
4. Electrostatic Precipitator Efficiency
95 percent
A-4
-------
5. Stack Height
500 feet
6. Drives for Fans and Pumps
Electric
7. Heat Rate
10,000 Btu/kwh
8. Coal Consumption
0.7505 lb/kwh
9. Coal Firing Rate
375,250 lb/hr
187.63 tons/hr
10. Flue Gas Flow Rate
1,056,204 SCFM (70°F, 1 atm, wet basis)
1,564,355 CFM (325°F, 1 atm, wet basis)
11. Fly Ash Leaving Furnace in Flue Gas
28,106 lb/hr
14.053 tons/hr
12. Fly Ash in Flue Gas Leaving Precipitator
Concentration = 0.1553 grain/SCF (70°F, 1 atm, wet basis)
Mass emission rate = 1405.3 lb/hr
13. Sulfur Oxide Emissions
502 = 22,200 lb/hr
= 11.100 tons/hr
= 266.40 tons/day
503 = 391.01 lb/hr
= 9384 lb/day
A-5
-------
Model B Plant
1. Location
Cairo, Illinois
2. Generating Capacity
1000 MW
3. Gas Cleaning System
No temperature limits apply. Unless some other gas
cleaning system is specified as an integral part of the
sulfur oxide control process, the power plant will be
assumed to be equipped with an electrostatic precipitator
operating on flue gas at a temperature of 325°F and having
an efficiency of 99.5 percent.
4. Stack Height
The plant will be assumed to have an 800-foot stack when
not equipped with a sulfur dioxide control system, and a
500-foot stack when equipped with a sulfur dioxide control
system of 90-percent efficiency.
5. Drives for Fans and Pumps
Either electric or steam turbine
6. Heat Rate
9600 Btu/kwh
7. Coal Consumption
0.7204 lb/kwh
8. Coal Firing Rate
720,400 lb/hr
360.20 tons/hr
A-6
-------
9. Flue Gas Flow Rate
2,027,690 SCFM(70°F, 1 atm, wet basis)
3,003,230 CFM (325°F, 1 atm, wet basis)
5,203,090 CFM (900°F, 1 atm, wet basis)
10. Fly Ash Leaving Furnace in Flue Gas
53,958 lb/hr
26.979 tons/hr
11. Fly Ash in Flue Gas Leaving Precipitator of 99.5 Percent
Efficiency
Concentration = 0.01553 grain/SCF (70°F, 1 atm, wet basis)
Mass emission rate = 269.79 lb/hr
12. Sulfur Oxide Emissions
S02 = 42,619 lb/hr
= 21.309 tons/hr
= 511.43 tons/day
SO = 750.66 lb/hr
o
= 18,016 lb/day
A-7
-------
Appendix B
COST FACTORS USED IN MODEL STUDIES
In an actual case, the busbar cost of electricity and the cost of
steam generated at a power plant are, of course, related to the cost of
fuel and other plant operating costs as well as to the fixed costs, which
are in turn related to the capital investment and the plant load factor.
In practice, it appears that the complexity of cost accounting makes it
difficult for even a utility company to assign costs precisely for an
individual power plant in its system.
In the present study, reasonable costs were assumed for individual
items such as fuel, steam, and electrical power. No attempt was made to
reconcile these costs, but it was appreciated that the values adopted may
actually not be entirely consistent with one another.
A. Fuel Costs
Coal
The cost of coal was estimated from data in "steam-Electric Plant
Factors (1967)," published by the National Coal Association. The costs
(^/million Btu) for coal as burned in Pennsylvania power plants were
averaged for coals having heating values greater than 13,000 Btu/lb, which
had higher unit costs than the lower-grade coals. The average cost per
million Btu was applied to the hypothetical coal assumed in the present
study (heating value 13,325 Btu/lb) to give the following costs:
31.2£/million Btu
$8.30/ton
The average costs of coal used in Illinois, Indiana, Ohio, and
Pennsylvania were lower, but applied to coals having heating values in
the range of 10,000 to 12,000 Btu/lb. Presumably these coals had generally
higher ash contents than the hypothetical coal.
B-l
-------
Natural Gas
An estimate of the cost of natural gas was taken as the rounded
average of the average costs of natural gas burned at power plants in
Illinois, Indiana, Ohio, and Pennsylvania. The data were taken from
"Steam-Electric Plant Factors (1967)." The value adopted was:
30£/million Btu
The heating value of the gas was assumed to be 1,000 Btu/CF.
No. 2 Fuel Oil
The cost of No. 2 fuel oil in central Pennsylvania was estimated
at SRI to be:
$4.25/barrel
B.
Water
The costs of water for various uses were taken from lists of
standard factors drawn from the literature or were figures used for esti-
mating purposes at SRI.
Raw makeup water
Once-through cooling water (temperature
80°F)
Process water
Boiler feed water
Retreatment of process steam
condensates for boiler use
Cost
2
-------
It was assumed that low-pressure steam was available as needed,
obtained by extraction from the power plant turbines. The value of low-
pressure steam extracted from turbines, or exhausted from a backpressure
turbine, was assumed to be proportional to the value of the high-pressure
steam in the ratio of the enthalpies of the steam at the two conditions.
D. Electrical Power
The busbar cost of electrical power generated and used in the power
plant itself (including the emission control system), after inspection of
data published by the Federal Power Commission and after discussions with
staff of the Pacific Gas & Electric Co., was assumed to be as follows:
Generating cost 3.2 mills/kwh
Fixed charges 2.3 mills/kwh
Total cost 5.5 mills/kwh
E. Salaries and Payroll Benefits
Salaries
Operating labor $3.75/hour
Supervision $4.75/hour
Payroll benefits
25 percent of labor and supervision
F. Maintenance
Taken as an appropriate percentage of capital investment to cover
both labor and materials.
G. Indirect Costs
Overhead
50 percent of labor, supervision, and maintenance
B-3
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H. Fixed Charges
The total fixed charge was adopted from the Tennessee Valley
Authority's Conceptual Design and Cost Study No. 2 (1969) for National
Air Pollution Control Administration, "Sulfur Oxide Removal from Power
Plant Stack Gas—Use of Limestone in Wet-Scrubbing Process." The break-
down of fixed charges (adapted from Federal Power Commission guidelines)
is given as follows:
Annual Percent of
Capital Investment
Depreciation (20-year life)
5.00
Insurance
0.25
Cost of capital (capital structure
assumed to be 50% debt and 50%
equity)
Bonds at 6% of average
undepreciated investment
1.50
Equity at 11% of average
undepreciated investment
2.75
Taxes
Federal
2.80
State
2.20
Total fixed charges
14.50
B-4
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Appendix C
STACKS FOR USE ON CONTROLLED
SULFUR DIOXIDE EMISSION SOURCES
Evidently, no clear policy has yet been developed with respect to
choice of the heights of stacks to be used on power plants equipped with
sulfur dioxide emission controls. The use of a high stack on an uncon-
trolled emission source reduces the maximum ground-level concentration
of sulfur dioxide at the point where the plume reaches the ground, but
it does not, of course, reduce the total emission of sulfur dioxide that
may affect areas beyond the range of influence of the stack. If emission
controls are used as well, not only is the maximum ground-level concen-
tration further reduced, but the quantity of sulfur dioxide affecting
the area beyond the range of influence of the stack is reduced also.
If, on the other hand, the use of emission controls is accompanied
by a reduction of the stack height, the reduction of the maximum ground-
level concentration of sulfur dioxide may be partly lost, depending upon
the balance between the reduction in stack height and the reduction in
emission. The point of maximum ground-level concentration will be
brought closer to the emission source.
There is reluctance to make major reductions in stack height, even
with installation of sulfur dioxide emission controls, because of concern
with other contaminants that may still remain in the flue gas and with
C4
the possibility of failures of the emission control system.
If the sulfur dioxide emission control system reduces the temper-
ature of the flue gas, the effect is to reduce the effective stack
height because the density of the gas is increased and hence its buoyancy
C3 C4
is reduced. ' When the hot gas is scrubbed, the effect of gas cool-
ing is counteracted to some degree by the lowering of the gas density due
to the increase in the water vapor content of the gas stream.
C-l
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Particular concern has been expressed about "negative buoyancy."
If the plume emerging from the stack is saturated and contains water
droplets, the subsequent evaporation of the drops in the atmosphere may
Cl C2
cool the gas sufficiently to cause the plume to descend rapidly. '
The presence of the droplets in the plume may directly produce such a
descent by increasing the effective density of the gas in the plume.
However, the author believes that the presence in stack gas of any such
relatively gross quantities of liquid would most likely be due to
entrainment from the scrubber, which can be prevented. It has been
Cl
suggested that at least some of the reported instances of rapid descent
of plumes of scrubbed gas may have resulted from downwash produced by
adjacent structures. Plumes of scrubbed gas that is saturated but still
substantially above ambient temperature have been commonly observed to
rise and disperse in the same general manner as do warm, dry plumes.
Careful observations of such plume behavior have been reported by Boyer
4- ! C1
et al.
The necessity for reheating scrubbed stack gases to ensure adequate
dispersion is not entirely clear, at least if the efficiency of the
scrubber is high and some minimum stack height is used. Using a hypo-
thetical 200-MW power plant unit with a 300-foot stack as an example,
C4
TVA calculated the effects of cooling (by scrubbing) and of reheating
the flue gas on the maximum ground-level concentration of sulfur dioxide,
assuming varying scrubber efficiencies. The results indicated that if
the scrubber had an efficiency of at least 90 percent and the stack
height was at least 300 feet, failure to reheat the gas would have only
a small effect upon the percentage reduction in the maximum ground-level
concentration of sulfur dioxide achieved by scrubbing. Applying a
commonly used rule of thumb that a stack should be at least 2.5 times
the height of adjacent structures, a typical power plant should in any
case use a stack at least 300 feet in height in order to avoid possible
downwash of the gas.
In the present study, it was assumed that in the Model A plant, the
treated flue gas would be discharged through the existing 500-foot stack.
C-2
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In the case of the Model B plant, it was assumed that the plant would be
equipped with an 800-foot stack if it had no sulfur dioxide emission
controls, and with a 500-foot stack if it were equipped with an emission
control system of 90 percent efficiency. Such a reduction in stack
height does not appear to be unreasonable. In the light of present
knowledge, a minimum stack height of 500 feet for the controlled emission
source seems appropriate.
The Model B control systems (both Cat-Ox and Wellman-Lord) were each
given credit for the differential between the costs of the 800- and 500-
foot stacks. A rough estimate of the cost differential for stacks 30
feet in inside diameter ($850,000) was made from Fig. 6-1 of "Control
I(C 2
Techniques for Sulfur Oxide Air Pollutants.'
REFERENCES
CI. Boyer, A.E., F. B. Kaylor, T. V. Ward, and F. J. Gottlich,
Atmospheric Dispersion of Saturated Stack Plumes, Des. Oper.
Air Pollut. Contr., Pap. MECAR Symp. 1968 (Pub. 1969), 32-40.
C2. National Air Pollution Control Administration, Control
Techniques for Sulfur Oxide Air Pollutants, Publication No.
AP-52 (Jan. 1969)
C3. Scorer, R. S., "Air Pollution," Pergamon Press, New York (1968)
C4. Tennessee Valley Authority, Sulfur Oxide Removal from Power
Plant Stack Gas—Use of Limestone in Wet-Scrubbing Process,
Conceptual Design and Cost Study No. 2 for National Air
Pollution Control Administration (1969)
C-3
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Appendix D
ESTIMATION OF THE VALUES OF SULFUR BY-PRODUCTS*
As applied to power plant flue gases, the Monsanto Cat-Ox system
can produce only 78-percent sulfuric acid. The Wellman-Lord system
produces concentrated sulfur dioxide, which has relatively insignificant
markets as such but can be converted to elemental sulfur or to either
93-percent or 98-percent sulfuric acid.
The estimates of the sulfur by-product prices at the two hypothetical
power plant sites (Altoona, Pennsylvania for the Model A plant, and Cairo,
Illinois for the Model B plant) were based on an assumed Gulf Coast price
for sulfur of $30/long ton. This Gulf Coast price was estimated during
the previous study0* as a likely average over the period to 1975.
Currently, sulfur prices have become chaotic as sulfur supplies have
shifted abruptly from shortages to surpluses. Apparently it is scarcely
possible to define an "established" price for sulfur. The development of
sulfur surpluses has followed a continuing period of low activity in the
market for fertilizers, which provides 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/iong 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 recognized, the
assumption of the $30/long ton price is probably as good as any that can
be made at this time. Prices at other locations can be estimated by
adding to the Gulf Coast price the cost of transportation from the Gulf
*
This material was prepared by F. Alan Ferguson, Industrial Economist, SRI.
D-l
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Coast. This procedure assumes, of course, that the incursion of the
sulfur from a new source will not be large enough to disrupt the local
markets.
Model A Plant - Altoona, Pennsylvania
The closest and most ready markets for sulfur by-products produced
at the hypothetical power plant Model A would be located in the Pittsburgh
area, about 120 miles away. The acid users in the Pittsburgh region cur-
rently consume over 500,000 tons of acid per year, a large portion of
which is used to clean (pickle) iron and steel products and to produce
ammonium sulfate from the ammonia in coke oven gas. Both of these uses
could employ 78-percent sulfuric acid as well as 93-percent or 98-percent
acid, and the total market is more than large enough to absorb all the
acid that could be produced at the Altoona plant. The Pittsburgh area
could also absorb elemental sulfur that might alternatively be produced
at Altoona.
The prices that could be obtained for the various possible by-
products producible at Altoona will depend upon a number of factors, a
few of which can be allowed for at the present time. On the basis of a
Gulf Coast sulfur price of $30/long ton, the delivered price of elemental
sulfur in the Pittsburgh area will be about &36.50/long ton. Since the
elemental sulfur now being sold in the Pittsburgh area is merchant sulfur,
that produced at Altoona could probably capture about half the existing
market if sold at the same price of $36.50/long ton. Since the cost of
transportation from Altoona to Pittsburgh would be about $3.50/long ton,
the f.o.b. price of the sulfur at Altoona might be $33/long ton.
Estimating the price of Altoona sulfuric acid is more difficult
because some of the Pittsburgh area acid is made for captive consumption
and the rest is made and sold as merchant acid. The prices of captive and
merchant acid can differ by as much as a factor of two. The price of
merchant acid will vary according to whether the acid is sold on long
term contracts or when and as needed (spot sales). However, relatively
little acid is sold at the higher spot prices. The highest price the
plant at Altoona could expect to receive for its acid at Pittsburgh
D-2
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would be the same as that which would return a reasonable profit to an
acid producer operating a contact acid plant and using Gulf Coast sulfur
as raw material. The cost of producing acid from sulfur costing $36.50/
long ton would be about $13/short ton (100 percent acid basis). With the
allowance of about 15 percent as a return on investment, the corresponding
price of acid would be $15/short ton (100 percent acid basis). The
corresponding f.o.b. prices at Altoona can be estimated by subtracting
the transportation charges of $4/ton (100 percent acid basis) for
93-percent acid and $5/ton (100 percent acid basis) for 78-percent acid,
giving:
93-percent acid: $12/ton of 100 percent acid
78-percent acid: $ll/ton of 100 percent acid
If it should be necessary to displace existing captive producers to
capture adequate markets, the Altoona acid might have to sell at Pittsburgh
for about $2/ton less than the cost of production from sulfur, or $ll/ton.
The corresponding f.o.b. prices at Altoona would then be:
93-percent acid: $7/ton of 100-percent acid
78-percent acid: $6/ton of 100-percent acid
A detailed market analysis would be required to determine whether the
price of the Altoona acid would be nearer the upper end ($11 to $12/ton)
or the lower end ($6 to $7/ton) of the price range estimated.
Model B Plant - Cairo, Illinois
Finding markets for elemental sulfur produced at a plant at Cairo
should not be difficult. In 1965, over 1.2 million tons of elemental
sulfur were barged along the Ohio River and the upper Mississippi River.
Assuming a Gulf Coast sulfur price of $30/long ton, sulfur produced along
the rivers should sell for $30 plus the cost of barging sulfur from the
Gulf Coast to the actual plant site. The cost of barging elemental sulfur
from the Gulf to Cairo, Illinois is about $2.90/long ton. Hence, sulfur
produced at the Cairo plant should sell for about $33/long ton.
D-3
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If sulfuric acid, rather than elemental sulfur, should be produced,
finding markets for large quantities of acid might be substantially more
difficult. In 1966, only about 130,000 tons of acid were barged along
the Ohio River and 30,000 tons along the upper Mississippi River. From
the summary (Table D-l) of uses of sulfuric acid near three sites on
these two rivers, it is apparent that the only end use large enough to
consume the quantities of acid that could be supplied from the Cairo
plant would be the production of phosphate fertilizers.
A recovery system producing sulfuric acid at Cairo might supply
either existing phosphate fertilizer manufacturers or a new plant located
at or near Cairo to take advantage of the local source of acid. Supplying
such a new fertilizer manufacturer would probably yield the highest price
for the acid, although a more detailed market analysis would be required
to determine whether this would actually be the case. Nevertheless, the
sulfuric acid would have to be priced to provide the new fertilizer
producer with a competitive advantage over the existing suppliers. The
existing producers of phosphate fertilizers that are capable of supplying
fertilizer to the midwestern market at the lowest prices are probably
those who manufacture their product at plant sites in Louisiana and ship
it to the markets by barge. The price of the by-product acid at Cairo
must therefore permit the new producer to make his fertilizer for less
than the Louisiana producers can make and ship theirs.
In Table D-2 a comparison is made of the costs of producing 54-
percent wet process phosphoric acid in Louisiana and at the hypothetical
Cairo site. The data indicate that the Cairo producer could pay about
$32.20 for the amount of 100-percent sulfuric acid necessary to make 1.85
tons of 54-percent (Po0_) phosphoric acid, which is equivalent to 1.0 ton
2 5
of P 0_. Since 2.7 tons of sulfuric acid (100-percent basis) are required
Z 5
to produce the 1.85 tons of phosphoric acid, the hypothetical producer
could pay $11.94/ton (100-percent acid basis) for the sulfuric acid and
be able to make phosphoric acid to sell for the same price as that made
in Louisiana and shipped to the Cairo area. However, the sulfuric acid
would probably have to be offered at somewhat less than 11.94/ton in
order to provide incentive to the new phosphoric acid producer. Under
D-4
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these circumstances, the sulfuric acid could probably be sold for $10
to $ll/ton (100-percent acid basis). Since little or no transportation
would be required, 78-percent and 93-percent acid should sell for the
same price per unit quantity of 100-percent acid.
REFERENCES
Dl. 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.
D-5
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Table D-l
ESTIMATED SULFURIC ACID DEMAND IN SELECTED
PRODUCING AREAS, 1966
(thousands of short tons 100% HgSC^)
Cairo-Paducah St. Louis Chicago
Phosphate fertilizers 107 789 1,375
Ammonium sulfate 13 5 113
Iron and steel pickling — 3 169
Rayon and cellulose films and fibers — — 74
Petroleum refineries 10 81 78
Sulfonated detergents 15 57 33
Ammonium sulfate — 28 12
Titanium dioxide — 410
Hydrofluoric acid 45 — 38
Total Major End Uses 190 1,373 1,892
Source: A. D. Little, Inc.
D-6
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Table D-2
PHOSPHORIC ACID PRODUCTION COSTS
(per ton 100% P2^5^
Plant Location
Louisiana Cairo, Illinois
Quantity Cost/Ton Quantity Cost/Ton
Required 100% P„0„ Required 100% P„0_
——— —5— ——— 2,—5—
Phosphate rock (72 bpl)
$6.00/ton + $3.25 frt 3.2 tons $29.60
$6.00/ton + $5.85 frt 3.2 tons $37.92
Elemental sulfur
$30/long ton + $1.50 frt 0.82 25.83
0 0
HgSO^ manufacture $3.40 2.7 9.18 0 0
per ton
(excluding cost of
sulfur but including
15% return on investment)
Phosphoric acid manufacture 14.00 14.00
(excluding raw material
cost and profit)
Barge freight $3.00/ton 1.85 5.55 0
(phosphoric acid from
Louisiana to Cairo, 111.)
Subtotal
H2S04 purchase price $11.94/ton
(to match cost of producing
phosphoric acid at Cairo, 111.
and in Louisiana)
Total
$84.16 $51.92
0 2.7 32.24
$84.16 $84.16
Source: SRI
D-7
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