COST NOMOGRAPHS OF
SELECTED SULFUR DIOXIDE
ABATEMENT METHODS
Hittman Associates, Inc.
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COST NOMOGRAPHS OF
SELECTED SULFUR DIOXIDE
ABATEMENT METHODS
HIT-508
January, 1972
Prepared Under
Contract No. EHSD-71-43
Environmental Protection Agency
Office of Air Programs
HITTMAN ASSOCIATES, INC.
COLUMBIA, MARYLAND
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iii
LEGAL NOTICE
This report was prepared as an account of Government sponsored work.
Neither the United States, nor the Environmental Protection Agency, Office of
Air Programs (EP A-OAP). nor any person acting on behalf of EP A-OAP:
A. Makes any warranty or representation, expressed or implied with
respect to the accuracy, completeness, or usefulness of the information con-
tained in this report, or that the use of any information, apparatus, method,
or process disclosed in this report may not infringe privately owned rights;
or
B. Assumes any liabilities with respect to the use of, or for damages
resulting from the use of any information, apparatus, method, or process
disclosed in this report.
As used in the above, "person acting on behalf of EP A-OAP" includes
any employee or contractor of EP A-OAP, or employee of such contractor J to
the extent that such employee or contractor of EP A-OAP, or employee of such
contractor prepares, disseminates, or provides access to any information
pursuant to his employment or contract with EP A-OAP, or his employment
with such contractor.
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T ABLE OF CONTENTS
LEGAL NOTICE. . . .
T ABLE OF CONTENTS
. .. . .. .. ..
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LIST OF FIGURES.
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LIST OF TABLES. . .
I.
II.
III.
IV.
V.
VI.
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INTRODUCTION. .
NOMOGRAPHS. .
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PARAMETRIC APPLICATIONS. . . .
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DISCUSSION. . . . . . . . . . .
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B.
Basis of Process Cost Estimates.
Basic Methodology of Nomograph Preparation.
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A.
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C.
Discussion of Processes. . . . . .
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D.
Process Developmental Status. .
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E.
Desulfurization of Fuels. . .
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F.
Transportation Costs. .
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CONCLUSIONS AND RECOMMENDATIONS. .
REFERENCES. . . . . . . . . . . . .
.......
APPENDIX A - NOMOGRAPH EXAMPLES. .
APPENDIX B - COST ESTIMATES. . . . . .
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v
Page
ii
iii
iv
v
1-1
II-1
III - 1
IV-1
IV-1
IV-3
IV-6
IV - 43
IV - 50
IV - 64
V:1
VI-1
A-1
B-1
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Figure
No.
III - 1
IV-1
IV-2
IV-3
IV-4
IV-5
IV-6
IV-7.
IV-8
IV-9
IV - 1 0
IV - 11
IV -12
. IV - 13
IV - 14
IV - 15
IV - 16
LIST OF FIGURES
Title
Parametric Curves
Dry Limestone Injection Process
Wet Limestone Injection Process
Chemico MgO Scrubber System - Oil Fired Boilers
Babcock & Wilcox MgO Process
S02 Removal Efficiency for Venturi-Type Absorber
S02 Removal Efficiency for Floating Bed Absorber
Integrated Catalytic Oxidation System for New
Power Plant
Catalytic Oxidation Reheat System for Existing
Power Plant
Tyco Catalytic Chamber Process Flow Diagram
Molten Carbonate Process Flow Diagram
Absorber Operating Limits
Effect of S02-M2C03 Ratio Upon the Percent Removal
of S02
Estimated Remaining Coal Reserves of the United States
States on January 1, 1965
Effect of Specific Gravity of Separation on Coal
Washing Parameters
Sulfur Distribution of Remaining Bituminous Coal
Reserves (1965) .
Incremental Desulfurization Costs Based on Five
Reference Bituminous Coal Mines (1965 Dollars)
vii
Page
III - 3
IV-8
IV -13
IV - 1 7
IV - 19
IV - 21
IV - 22
IV - 25
IV - 27
IV - 30
IV - 34
IV - 37
IV - 39
IV - 58
IV - 59
IV - 61
IV - 63
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LIST OF TABLES
Table
No.
Title
II-l
Preliminary Computations
II-2
II- 3
Process Computations
Ranges of Input Parameters
IV-l
Nomograph Parameters Based on the Kellogg
Assumptions
IV-2
S02 Removal Processes - Base Case
Carbonate Makeup Requirements Assuming 88 Percent
Li2C03 Recovery
IV-3
IV-4
IV-5
Developmental Classification of S02 Removal Processes
Domestic Desulfurization Data for United States PAD
Districts
ix
Pag-e
II-2
II-3
II-5
IV-2
IV-5
IV-41
IV-44
IV - 54
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1-1
1.
INTRODUCTION
The continued economic growth of the United States requires ever-
increasing supplies of electrical power. From past events, the power industry
has needed to double its capacity each decade merely to remain even with the
demand. This growth trend is further expected to continue through at least
the end of this century. Nuclear power generating facilities will provide a
substantial percentage of the required power through the next 30 years.
However, fossil fuels will continue to be used in greater amounts until about
the turn of the century.
A growing awareness on the part of government, industry, and the
general population of the potential harmful effects on the environment and the
ecology from continued industrial growth has spurred a nationwide program
to control and to abate environmental impacts. As part of this program, large
efforts and many millions of dollars have been expended to find ways in which
the emissions of fossil fuel combustion from power generating facilities could
be minimized. A variety of techniques for the control of effluents from power
plants have been proposed, designed, or constructed.
In order to provide a basis for assessing the costs and effectiveness of
alternative sulfur dioxide control systems, the Environmental Protection
Agency (EP A) contracted with Hittman Associates, Inc., to perform a major
study to provide a set of nomographs which would allow rapid and accurate
assessment of these alternatives for control. This document contains the
results of that study.
Sulfur emissions can be controlled by either stack gas control processes,
desulfurization of oil or. coal, or the burning of naturally-occurring low sulfur
fuels. The economic aspects involved are the capital and operating costs of
stack gas and desulfurization facilities and the incremental costs of purchasing
and shipping low sulfur fuels.
The main emphasis was placed on the system and economic analysis of
six stack gas processes. The most current technical and economic data were
compiled and evaluated. The major cost parameters were identified and a
method whereby the interrelation of these parameters could be depicted
graphically was developed. The resulting nomographs allow the calculation
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1-2
of operating costs for any combination of plant size, fuel,
load factor, fixed charge rate, additive cost, etc. Thus,
process can be selected for any given power plant.
sulfur content,
the optim urn control
In addition to stack gas control methods, the economic feasibility of
coal and oil desulfurization was studied. The costs per ton of coal or per
barrel of oil were determined given a variety of conditions. Nomographs were
developed which allow one to determine desulfurization costs and to compare
these costs with those of the stack gas removal processes.
The cost of transporting bulk reactants or low sulfur fuels can be signi-
ficant when studying control economics. For this reason, a nomograph was
developed which allows the determination of transportation costs. All fuel
forms (coal, gas, and oil) and modes of transportation were considered.
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II-1
II. NOMOGRAPH8
The major purpose of the nomographs presented in this section is the
calculation of 802 abatement method costs. The nomographs dealing with
stack gas 802 removal systems calculate the fixed and variable costs of the
system when operated under known conditions. Typical inputs are the plant
size, load factor, heat rate, fuel, and the sulfur content in fuel. Annual costs,
as well as cost per kilowatt-hour, can be obtained.
Nomographs appearing in this section include the following: (Page
numbers in parentheses refer to the separate nomograph section at the end
of this chapter.)
.
Plant size and Fuel Consumption Nomographs (pages 1-4)
Dry Limestone Nomographs (pages 5-9)
Wet Limestone Nomographs (pages 10-13)
Magnesium Oxide Nomographs (pages 14-16)
Catalytic Oxide Nomographs (pages 17-19)
Modified Chamber Nomographs (pages 20-22)
Molten Carbonate Nomographs (pages 23-26)
Operating Cost Nomographs (pages 27 -29)
Desulfurization and Transportation Nomographs (pages 30-35)
.
.
.
.
.
.
.
.
To allow convenient use of the process nomographs, work sheets have
been included. These sheets can be reproduced to provide any number of
additional forms. They provide the proper format for the step-by-step utili-
zation of the nomographs. Table II -1 refers to nomographs on pages 1-4 in
which the basic computations necessary for all processes are performed.
Table II-2 refers to the process nomographs on pages 5-26. Table II-3 is
included to provide a first estimate for all required input variables. A low,
typical, and high value for each parameter is included to bracket the expected
range. Appendix A tabulates example cases for each process. These
examples will assist the individual's first attempts at using the nomographs.
Utilizing the nomograph on page 27, the annual quantities of sulfur
removed and sulfur released upon combustion may be determined as a function
of energy utilization, sulfur content of the fuel and the process 802 removal
efficiency. The nomographs on pages 28 and 29 provide alignment charts
used to convert annual operating costs into mills per kilowatt-hour, dollars
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T ABLE II -1. INITIAL CALCULATIONS
H
H
I
tV
Page
I
t
o t
t
npu s u;pu s
1 I A. Heat Rate (Optional)
1. Plant overall electrical 2. Plant net heat rate, Btu/Kw-hr
! efficiency, %
!
1 I B. Capital Recovery Factor I
I 3. Interest rate % I I
I 4. Years amortized 5. Capital recovery factor, %/yr I
; 6. Recurring costs, %/yr 7. Fixed charge rate, %/yr
(add 5 and 6)
2 C. Energy Utilization
see 2 above plus. . .
8. Net generating capacity, Mw
9. Plant load factor, % 10. Annual energy production,
109 Kw-hr/yr
11. Energy utilization, 1012 Btu/yr
3 ID Fuel Cost and Consumption I
I s~e 11 above plus. . .
I 12. Heat value 13. Annual fuel consumption
I
! 14. Fuel cost (optional) 15. Energy cost (optional) 9/ 106 Btu
4 E. Sulfur Content Unit Conversion I
see 12 above plus. . . I 6
16. Weight percent of sulfur in fuel i 17. Sulfur content, Ib/10 Btu
i I
4 i F. Flue Gas Flow Rate 6
I 2 and 8 above 18.
! see Flue gas flow rate, 10 sefm
I ! -
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II- 3
TABLE II-2. PROCESS CALCULATIONS
Pal!e Inputs Outputs
Dry Limestone Process
5 A. Efficiency
19. Stoichiometric ratio 20. Sulfur dioxide removal efficiency, %
6 B. Limestone Use Factor
see 17 and 19 above plus. Limestone use factor /lb/ 106 Btu
21. Limestone purity, wt-% of carbonate 22.
7 C. Fixed Cost
see 7, 18, and 22 above. 23. Fixed capital cost, 106$
24. Fixed cost, 106$/yr
8 D. Othe r Operatinl! Costs
see 11 and 22 above. 25. Other operating costs, 106$/yr
8 E. Grindinl! Correction (Optional)
26. Limestone particle size, 11 27. Incremental cost of grinding, $/ton
9 F. Additive Cost
I
I see 11 and 22 above plus. . . 106$/yr
28. Total reactant cost, $/ton-add delivered 29. Annual additive cost,
limestone cost, incremental cost of
grinding (see 27), and disposal cost
Dry Limestone Total Operating Cost, 106$/yr Add 24, 25, and 29
Wet Limestone Process
10 A. Efficiency
Assumed constant at 85%
10 B. Limestone Use Factor
see 17 above plus.
30. Stoichiometric ratio Limestone use factor ,lb/ 106 Btu
31. Limestone purity, wt-'70 of carbonate 32.
11 C. Fixed cost
see 7, 18, and 32 above 33. Fixed capital c~st, 106$
34. Fixed cost, 10 $/yr
12 D. Other Operatinl! Costs
see 11 and 32 above 35. Other operating costs, 106$/yr
12 E. Reheat Correction (Optional)
see 18 above plus. " 0 37. Credit, 106$/yr
36. Reheat temperature, F
13 F. Additive Cost
see 11 and 32 above plus. . . 106$/yr
38. Total reactant cost, $/ton-add delivered 39. Annual additive cost,
limestone cost and disposal cost
Wet Limestone Total Operating Cost, 106$/yr Add 34, 35, and 39, subtract 37
Magnesium Oxide Process
14 A. Fixed Cost
see 7, 17, and 18 40. Fixed capital cost, 106$
41. Fixed cost, 106$/yr
15 B. By-Product Acid Credit
3
see 11 and 17 above plus. 42. Sulfur removed~ 10 ton/yr
43. $/ton of 98.5,,/0 H2S04 44. Acid credit, 10 $/yr
16 C. Raw Material Cost
see 11 above plus. 106$/yr
45. $/ton of magnesia 46. Annual MgO cost,
I 16 D. Other Operatinl! Costs
I see 11 and 17 above 47. Other operating costs, 106$/yr
I
I Magnesium Oxide Total Operating Cost, 106$/yr Add 41, 46, and 47, subtract 44
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II-4
TABLE II-2. PROCESS CALCULATIONS (Continued)
Page Inputs Outputs
Catalytic Oxidation Process
17 A. Fixed Cost
see 7, 17, and 18 above 48. Fixed capital cost, 106$
49. Fixed cost, 106$/ yr
18 B. By-Product Acid Credit
see 17 and 18 above plus. . 50. 3
Sulfur removedt; 10 tons/yr
51. $/ton of acid 52. Acid credit, 10 $/yr
19 C. Raw Material Cost
see 11 above plus. . . Quantity of V205 catalyst, ft3/~r
53. % ash in fuel 3 54.
55. V205 cost ($/ft ) 56. Annual V 205 catalyst cost, 10 $/yr
19 D. Other Operating Costs
see 11 above 57. Other operating costs, 6
10 $/yr
Catalytic Oxidation Total Operating Cost, 6 Add 49, 56, and 57, subtract 52
10 $/yr
TYCO Modified Chamber Process
20 A. Fixed Cost
see 7 and 18 above 58. . Fixed capital cost, 106$
59. Fixed cost, 106$/yr
21 B. Raw Material Cost
see 11 above plus. . . 3 6
60. Charcoal catalyst cost, $/ ft 61. Annual catalyst cost, 10 $/yr
21 C. Other Operating Costs
see 11 above 62. Other operating costs, 106$/yr
22 D. By- Product Acid Credits
see 11 and 17 3
63. Sulfur removedj 10 ton/yr
64. 80"/0 H2S04, 10 ton/yr
65. $/ton of 80"/0 H2S04 66. By- product H~S04' 106$/yr
67. NOx volume content, ppm 68. 600/0 HN03, 10 ton/yr
69. $/ton of 60"/0 HN03 VO. By-product HN03' 106$/yr
TYCO Modified Chamber Process, 106$/yr Add 59,61, and 62, subtract 66 and 70
Molten Carbonate Process
23 A. Fixed Cost
see 7, 11, and 17 above 71. Fixed capital cost, 106$
72. Fixed cost, 106$/yr
24 B. Other Operating Costs
see 11 and 17 above 73. Other operating costs, 6
10 $/yr
24 C. By-Product Sulfur Credit
3
see 11 and 17 above plus. . . 74. Sulfur removed, 10 ton/yr 6
75. $/ ton of sulfur 76. By-product sulfur credit, 10 $/yr
25 D. Raw Material Costs
& see 11 and 74 above plus. .
26 6
77. $/ ton of coke 78. Annual coke cost, 10 $/yr 6 (1)
79. Percent of ash in fuel 80. M2C03 Replacement cost, 10 $/yr (2)
81. M2C03 Replacement cost, 106$/yr (3)
82. Percent of chlorine in fuel 83. M2C03 Replacement cost, 106$/yr (4)
Molten Carbonate Total Operating Cost, 6 Add 72,73,78,80,81, and 83,
10 $/yr
subtract 76
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II- 5
TABLE II-3. RANGES OF INPUT PARAMETERS
Page
Nomograph No. Input Identification Low Typical High
Initial 1 1 20 30 40
Calculations 3 3 6 9
4 5 10-15 30
6 1 2-3 4
2 8 100 500 1,000
9 20 50 75
3 12 (coal only) 6000 12,000 14,000
14 (coal only, $10/ ton) 0.3 O. 6 1.0
4 16 (coal only) 2 3.5 5.0
Dry 5 19 1 2 3
Limestone 6 21 80 90 95
9 28 Limestone 2 4 7
Grinding (size dependent - see output 27)
Disposal 0.4 0.5 1.0
Total 2.4 4.5 8.0
Wet 10 30 1 1.3 1.5
Limestone 31 80 90 95
12 36 150 250 275
13 38 (see 28)
Magnesium 15 43 5 12 15
Oxide 16 45 40 50 80
Catalytic 18 51 1 5 10
Oxidation 19 53 1 10 15
55 30 38 50
TYCO 21 60 10 14 20
Modified 22 65 1 5 10
Chamber 67 250 400 ; 1,200
69 10 15 35
Molten 24 75 5 20 25
Carbonate 25 77 5 8 20
79 1 10 15
26 82 0.01 0.05 0.5
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II-6
per unit of fuel consumption, and dollars per ton of sulfur removed.
Pages 30-32 calculate the cost of fuel desulfurization. Examples of the
use of these nomographs, as well as a technical discussion of fuel desulfur-
ization, can be found in Section IV.E. Pages 33-35 include nomographs which
estimate the cost of transporting oil, gas, coal, and solid reactants.
A cost sensitivity analysis was performed utilizing the nomographs.
Major parameters such as plant size, load factor, sulfur content, and addi-
tive costs were varied and the effect on total operating cost determined. The
results of the analysis are presented in Section III.
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NOTICE
The nomographs which follow are based on the best currently available
economic data. It is expected, however, that these data will change as pro-
cess development progresses. Prototype tests currently being conducted or
in the planning stages are expected to make significant contributions to the
economic data base. Improvements in, or modifications to, system com-
ponents are also expected to affect economics.
Economic comparisons of processes are meaningful only if they are at
the same stage of development. Comparison of a fully developed process
with one not yet in the prototype stage may not be accurate since the economics
of the least developed process are based on assumptions which have not been
thoroughly tested. Section IV. D discusses development status of the systems.
Many secondary economic variables have been held constant at typical
or maximum values. This was required to simplify the nomographs. Sec-
tion IV. C of this report discusses these variables in detail.
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-------
III - 1
III.
PARAMETRIC APPLICATIONS
One of the more revealing applications of the process nomograph system
is in preparing parametric curves. Because of the ease of manipulating the
basic economic variables, one can perform a parametric study in a relatively
short period of time. Six such parametric studies were made to exemplify the
kind of comparative analysis the nomographs make possible. The cost of S02
control in terms of mills/Kw-hr were calculated for over a dozen cases. Each
case was run through each of the six process nomographs. The method involved
varying one parameter at a time, keeping all others constant. The cost impact
of the varying parameter was thus revealed. In Figure III-I., curves A-F sum-
marize the parametric results. Each curve is based on the same basic set of
parameters which are summarized below:
Base Case Inputs
Net generating capacity
Sulfur content
Load factor
Additive costs
500 Mw
3.5 wt-%
75%
Typical (see Table II - 3)
Typical (see Table II - 3)
20%/yr
11,000 Btu/lb (coal)
By-product revenue
Fixed charge rate
Plant heat rate
The base case parameters were varied as follows:
A. Net generating capacity 200, 500, 1000 Mw
B. Sulfur content 1, 3.5, 5 wt-%
C. Load factor 20, 50, 75
D. Additive costs Low, typical, high
(see Table II - 3)
E. By-product revenue Low, typical,high
(see Table II - 3)
F. Fixed charge rate 5, 10, 20 %/yr
-------
III - 2
These input variations were entered into each of the six process nomo-
graphs. The numbered curves appearing in Figure III-1 correspond to one of
the six processes as follows:
1.
2.
3.
4.
5.
6.
Dry limestone
Wet limestone
Magnesium oxide
TYCO modified chamber
Catalytic oxide
Molten carbonate
In studying the resulting variations in control costs, many significant
conclusions can be drawn. For instance, in Curve A, note the significant
increase in cost of Processes 3-6 as plant size is decreased. Process 2,
however, is less affected over the range. Note also that costs of Processes
2, 4, and 5 approach one another at the 1000 Mw level. When sulfur content
was varied, costs were affected as shown in Curve B. Note that the limestone
processes are affected most adversely. The other processes having some
sulfur by-product are generally affected less. Note that Processes 3 and 5
actually decrease in cost as sulfur content is increased. Curve C shows the
dramatic change in cost when load factor is varied. As is expected, the lime-
stone processes are least affected since they are the least capital intensive
processes. Similar behavior patterns are evident when the cost of bulk
reactants, fixed charge rate, and the value of by-products are varied (Curves
. D-F). Thus, the nomographs provide a flexible and easy-to-use method of
analyzing any of the important economic parameters.
-------
H
.J:1
I
!:=
.~
-
rJ)
.-I
......
.,...j
S.
..
.-I
o
H
...
s::
o
U
C"J8
o
U'J.
.....
o
...
rJ)
o
U
......
ro
...
s::
Q)
S
Q)
H
CJ
s::
~
2
1
o
Net Generating Capacity,
Mw
6
4
2
o
C.
0/0
Load Factor,
Figure lIT-I.
B.
Weight Percent of Sulfur Fuel
o
D.
Typical High
Cost of Bulk Reactants
~
~
......
I
CI.:)
Low
Parametric Curves
-------
r-. ~ 2
...c= ...c
I I
~ ~
~ ..!;:::
- -
to U)
......
- :::i
..... 1 1
~ .....
~
o
10
E.
Fixed Charge Rate,
%/yr
H
H
H
I
~
3
o
Low
Typical
High
F.
By- Product Value
Figure III-I.
Parametric Curves (Continued)
-------
IV-1
IV.
DISCUSSION
A.
Basic Methodology of Nomograph Preparation
The overall purpose of the nomographs is to present fixed capital and
annual operating costs for six S02 removal processes. These costs are pre-
sented as a function of the major independent variables such as plant size,
plant load factor, plant efficiency, and the weight-percent of sulfur in the fuel.
Fixed capital and annual operating costs were based primarily upon the M. W.
Kellogg Company unified cost estimates. These estimates were for plant
capacities and coal sulfur levels as shown in Table IV -1. Other sources for
costs were employed such as the reports from the process developers, EPA,
TV A, and other consultants. Cost estimates were also performed by Hittman
Associates to extend plant size from 100 to 2000 Mw and sulfur content from
0.5 to 5 lb S/106 Btu.
A cost sensitivity analysis was performed in which each independent
parameter was ranked in terms of its significance to overall cost. If varying
the parameter from its minimum to its maximum possible value significantly
affected the total operating cost, it was introduced into the nomographs as a
variable. If it was not significant, a mean value was selected and it was
handled as a constant. For example, the variability of operating labor rates
always resulted in only a small overall effect on operating cost and therefore
was fixed at the average rate of $4. 50/hr. This analysis was necessary to
minimize the number of variables for practical application.
The fixed capital costs for each process are presented as a function of
the volumetric flow rate of the flue gas into the system in cubic feet per minute
(scfm) at standard conditions. Flue gas flow rate is determined as a function
of plant capacity (Mw) and plant heat rate (Btu/Kw-hr) in auxiliary nomograph
No. 1. In most of the processes, fixed capital costs also depend upon the sul-
fur content of the fuel. For coal, sulfur content is determined from the
following equation:
4
SC = weight-percent of sulfur x 10 , lb S/106 Btu
heating value of the coal
Sulfur content is plotted for coal and oil on nomograph page 4.
-------
IV-2
TABLE IV-1. NOMOGRAPH PARAMETERS
BASED ON THE KELLOGG ASSUMPTIONS
Plant Wt- % Sulfu r Energy Flue Gas
Capacity of Sulfur Content Utilization Flow Rate
lb S/106 Btu 12 6
Case Mw in Coal 10 Btul yr 10 scfm
1 175 3. 5 2.92 11 .32
2 350 2. 1. 67 22 .64
3 350 3.5 2.92 22 .64
4 350 5. 4.17 22 .64
5 400 3. 5 2.92 25 .72
6 700 3.5 2.92 44 1. 27
7 1000 2. 1. 67 63 1.8
8 1000 3. 5 2.92 63 1.8
9 1000 5. 4.17 63 1.8
NOTES:
Plant Heat Rate = 9000 Btu/kw-hr
Overall Plant Efficiency = 38%
Heating Value of Coal = 11,980 Btu/lb as received
Plant Load Factor = 80% (70100 hr/yr)
Annual Fuel Consumption for Base Case = 2.63 x 106 tons/yr
Case s 1 - 4, Dry Lime stone Froce s s only
Cases 5- 9, All Other Processes
Case 8, Base Case
-------
IV-3
In the nomographs, the total annual operating cost is determined for a
given condition by summing the fixed cost, the raw materials costs, and the
other operating costs, and by subtracting the credit for by-products resulting
from the process. The fixed cost in dollars per year depends upon the capital
cost and the fixed charge rate. The fixed charge rate in percent per year is a
function of the interest rate, amortization period, and the recurring cost per-
cent. It is determined by adding the recurring cost percentage to the capital
recovery factor obtained from nomograph page 1. Other operating costs in-
clude utilities, direct labor, supervision, maintenance, overheads, and waste
disposal. Other operating cost magnitudes also account for other credits that
are unique for each process such as stack gas cooling and boiler feedwater
heating.
The annual operating costs and credits in the process nomographs depend
upon plant capacity through the parameter, energy utilization. Energy utiliza-
tion is equal to the product of the annual fuel consumption (at a given plant load
factor) and the heating value of the coal. It is determined from nomograph
page 2. These costs also depend upon the parameters: sulfur content of fuel,
ash content of the fuel, chloride content of the fuel, NO content of the flue gas,
x
raw material unit costs, by-product unit costs, etc.
In addition to the Plant Size and Fuel Consumption Nomograph (pages 1-4)
and the six S02 removal process nomographs, two other nomographs were con-
structed. On pages 27-29, operating costs can be converted to other param-
eters for comparison purposes. The sulfur emitted from the stack is also
determined as a function of sulfur content, energy utilization, and" 802 removal
efficiency. On pages 30- 35, the costs of oil and coal desulfurization are com-
puted. Also, the costs of oil, coal, and gas transportation are determined.
B.
Basis of Process Cost Estimates
The basic approach to estimating cost involved the comparison between
a new plant with and without S02 control. Conventional equipment which is
eliminated due to the addition of the control equipment was credited to the
process capital cost. The increase or decrease in the cost of dust collectors,
induced draft fans, ductwork, insulation, etc., required when adding the con-
trol equipment was considered a debit or credit to the control process.
-------
IV-4
Cost estimates have been performed by EP A, Hittman Associates, M.
W. Kellogg Company, TV A, and Singmaster & Breyer, along with the process
developers. In general, these estimates were difficult to compare since no
single set of cost estimating techniques were used. It was found that com-
parisons between estimates could be very misleading. A large amount of
effort was sometimes necessary in order to discover the reasons for various
discrepancies in the data. These difficulties point to the need for a uniform
cost guideline. Such a guideline would allow comparisons between estimates,
would allow an individual to adjust an estimate based on his own needs and
would eliminate many potential errors and omissions.
Primarily because estimates from different organizations were difficult
or impossible to compare on a single basis, it was decided that the most uni-
form set of estimates would be utilized as a basis for the nomographs. It was
reasoned that the process-to-process comparisons would only be fair if a single
methodology was used. In reviewing all the estimating work completed in 1971,
it was found that the M. W. Kellogg Company had provided EP A with estimates
of the six stack gas processes (among others) based on uniform methods for
estimation and presentation. These estimates were for a base case of 1000 Mw
plant size and a 3.5 wt-% sulfur content in the fuel. These estimates were used
as the basis of the parametric analysis performed by Hittman Associates.
The breakdown of the capital cost estimates performed by the Kellogg
Company for the base case is given in Table IV - 2. The base case for the
. dry limestone process was a plant capacity of 350 Mw and a coal whose sulfur
content was 3.5 percent by weight. The base case for the remaining five S02
removal processes was a 1000 Mw plant burning the same coal. The costs of
induced draft fans and dust collection equipment required with the process are
included in the major equipment cost. Credit for the cost of these components
sized for a plant with no controls are given at the bottom of the table. A con-
tractor's fee was included under "Other Cos ts." The fixed capital cost or net
cost investment does not include any contingency or interest during construction.
The annual operating costs for the base case of each process are pre-
sented in Tables B-1 through B- 6 of Appendix B. ' These estimates and the
fixed costs of Table IV - 2 were transmitted by the Kellogg Company to Hittman
Associates in Reference 1. The operating cost assumptions are indicated in
-------
TABLE IV-2. S02 REI'v10V AL PROCESSES - BASE CASE IV-5
(Costs in $103)
Cfit-Ox Dry
Molten Wet New or Magnesia Limestone
Carbonate Limestone Retrofit Tyco Case 350 Mw
Major Equipment Cost:
Direct materials 3,637 5,395 13,525c 9,754 8,422 664
Subcontracts 4,231 324 1,360 6,949d 1,665 91
Ductwork, Duct Valves, and 3,565a 4,231a
Expansion Joints 937 811 1,182 124
Piping, Electrical, and Instruments 2,830 1,323 1,940 2,360 2,173 166
Foundations, Concrete and Steel
Structures, Buildings, Paint, and
In~ulation:
Direct materials 474 340 440 410 470 76
Subcontracts 673 170 340 290 290 70
Subtotal - Materials and Subcontracts 15,410 8,'489 21,836 20,574 14,202 1, 191
Other Costs:
Construction forces, home office
engineering, procurement, insur-
ance, startup, and contractor's fee 14,490 6,711 17,064 10,226 11, 398 1,409
Subtotal Investment Cost for Process
Plant 29,900 15,200 38,900 30,800 25,600 2,600
Other Process Units Investment Costs:
Claus sulfur plant 1,600
Sulfuric acid plant including H2S04 1,400b 1,300b 5,600
storage
Electrostatic precipitatoJ:1 7, 500f 7, 500r 2,400e
Deduct Standard Power Plant Induced
Draft Fans and Ducts (2,300) (2,300) (2,300) (2,300) (2,300) (700)
Deduct Standard Power Pl~nt Electro-
static Precipitatore (2,100) (2,100) (2,100) 0 (2,100) (900)
NET COST INVESTMENT 34,600 10,800 43,400 29,800 26,800 3,400
NOTES:
) = Deduction
aCosts also include duct insulation.
bCost is for sulfuric acid storage only.
~Cost includes V 205 catalyst.
dCost includes nitric acid storage.
e3000F electrostatic precipitator.
rDOOoF electrostatic precipitator.
-------
IV-6
the tables. Labor rates were assumed based on a plant location in Cincinnati,
Ohio. The stream time corresponds to a plant load factor of 80 percent. The
percentage of fixed capital investment accounting for maintenance labor and
materials was a variable. The unit cost for electric power under cost of
utilities was assumed to be $0.00675 /Kw-hr. Solid waste disposal was assumed
to cost $0. 50/ton. The fixed charge rate was assumed to be constant at 18.22
percent per year in these estimates. This parameter can be varied in the
nomographs when the fixed cost is required under various economic conditions.
The total operating cost is defined as being the net production cost wherein
account has been taken of the various credits.
Cost estimate data points were extended beyond the base case through
an additional effort performed by M. W. Kellogg Company. Reference 2 pre-
sents the effects on costs of varying the plant capacity and the sulfur content
of coal. Results of the parametric study are shown for each S02 removal
process in Tables B-7 through B-12 of Appendix B. One important assumption
in the study was that the concentration of sulfur dioxide emitted in the power
plant effluent gas be held constant for all cases at approximately 250 ppm. In
the tables, the percent of sulfur removed from the flue gas or the S02 removal
efficiency of each case differs somewhat, though not significantly, from that
of the base case.
The total gross production cost less the credits yields the total net pro-
duction cost for a given case. In determining credits, the unit costs for acid
. are for the concentrated acid. This magnitude is reduced proportionately to
obtain the unit cost at the lesser concentration.
C.
Discussion of Processes
1.
Dry Limestone Process
a. Process History. The first active program in the .United States
for 'development of a dry limestone: injection process was initiated in 1964 by
the Process Control Engineering Program of EP A. As part of this program,
a conceptual design study was conducted by the TV A for the development of a
large scale prototype. This prototype design was scheduled for testing in 1969
at TV A's Shawnee and Paradise plants in Kentucky.
-------
IV-7
Presently, the process is undergoing tests at the Shawnee plant near
Paducah, Kentucky (Ref. 3). Current demonstrations are achieving from 20
to 30 percent control of sulfur oxides. The EP A has indicated that an ultimate
goal of 50 percent efficiency is being sought. The current program is for
$3.3 million and 18 months (Ref. 4).
The process can use dolomite, which is a mineral containing
equal parts of calcium and magnesium carbonate (commonly symbolized by
MgC03 . CaC03)' or any alkaline earth such as limestone which, by definition,
contains more than 50 percent calcium carbonate (CaC03). Limestones are
abundant in most areas of the country, but their compositions vary widely.
Dolomite limestones are those containing an appreciable percentage of mag-
nesium carbonate (Mgc03).
b. Process Description. The process is initiated with the limestone
and/or dolomite being pulverized and fed into the high-temperature combustion
zone of the furnace where it is calcined (oxidized) to the active oxide forms
CaO and MgO. Figure IV -1 provides a schematic of the process. The reactions
are as follows:
CaC03(s)
MgC03(s)
> CaO(s) + C02(g)
>MgO(s) + C02(g)
The reaction of the pul verized- calcined additives with S02 and
oxygen at temperatures above 12000F forms gypsum or calcium sulfate
(CaS04) and magnesium sulfate (MgS04) which, along with the unreacted lime-
stone and/ or dolomite and the normal particulates (e. g., fly ash), are carried
out of the furnace in the flue gases. The calcium and/or magnesium sulfates
are removed with the fly ash in conventional dry scrubbers, either electo-
static precipitators or cyclone separators. Additional electrostatic precipi-
tator capacity may be required to maintain the particulate below regulated
values. The gypsum and other sulfates, along with the unreacted limestone,
are. handled as waste along with the fly ash. Additional disposal capacity will
be required to handle the increased waste. Because of the chemical compo-
sition of the resultant wastes, a water pollution problem may exist, depending
upon the disposal system design and local conditions. This is especially true
in the case of a sluice / settling pond system in which pond overflows due to
heavy rains may damage stream ecology. This problem must be handled on
a site-by-site basis.
-------
From
Yard
Storage
H
o
p
p
e
r
Coal
Furnace
A" !
Hea\:rl
Figure IV -1.
Dry Limestone Injection Process
I-<
<:
1
co
To
Stack
-. Dust Collector
-------
IV-9
c. Process Economics. Of all the processes, the dry limestone
process is the least complex. It requires a minimum amount of equipment
and is, therefore, the least capital intensive. It is also the process most
easily retrofitted on an existing plant. External receiving hopper, storage
facilities, pulverizer, and injection system plus conveyors between these units
are all that are necessary for limestone injection. A more efficient dust col-
lector and higher capacity disposal system make up the balance of the components.
The capital and variable costs of the system are displayed on nomo-
graph pages 5-9. Because the selection of the stoichiometric ratio (1. 0 stoichio-
metric refers to the amount of reactant required to eliminate all of the S02 as-
suming complete carbonate utilization) is known to significantly affect both S02
removal efficiency and operating cost, it has been introduced as a variable.
Generally, a range of between 1. 0 and 3.0 stoichiometry is selected. The se-
lection of values higher than 3.0 normally results in an unacceptable amount
of limestone entering the system. This material may overwhelm the dust col-
lection equipment and/ or result in excessive corrosion or erosion of boiler
int e rnals .
Graph A, which shows process efficiency versus stoichiometric ratio,
is a reasonable interpretation of currently available date (Refs. 5 through 8).
Efficiency varies widely because of the many contributing factors. Limestone
chemical and physical characteristics cause significant variations in sorption
efficiency. Smaller limestone particle size results in higher sorption efficiency.
The time -temperature profile of the injected limestone particle s has a signifi-
. cant effect on the calcination/ sorption processes. The distribution of the par-
ticles as they enter the fire box also contributes to the variability in S02 remov-
al efficiency. These factors and others are being studied in the attempt to in-
crease the process effectiveness.
Graph B calculates the limestone use factor (LUF). The stoichi-
ometry, the amount of sulfur in the fuel, and the limestone purity determine the
amount of limestone used per Btu of energy liberation. The factor is defined
below:
L UF = S x R x CS R
P
where:
LUF =
S =
limestone use factor, lb of limestone/106 Btu
sulfur content in fuel, lb of sulfur/106 Btu
-------
IV - 10
R
CSR
p
=
stoichiometric ratio
ratio of CaC03 to sulfur at 1. 0 stoichiometry, 3.12
limestone purity ~ percent of CaC03 in limestone, %/100
=
=
The LUF is utilized a number of times in obtaining cost data.
Because of the similarities in variables between the dry and wet limestone
processes, the LUF graph is used for both processes.
Graph C allows calculations of the capital and fixed costs. The
major variables affecting capital cost are the limestone use factor and plant
flow rate. Graph E allows for the adjustment of costs when the use of a lime-
stone particle size smaller than normal is found to be desirable. . The graph
should be used in conjunction with test data which shows a definite improve-
ment in efficiency with finer grinding. If no test data exist, it is best to over-
look the correction and assume standard grinding (80 percent of the material
ground to less than 44 microns, which is equivalent to passing a No. 325
screen) .
Graph F accounts for limestone usage. The nomograph calculates
the amount of limestone used per year and the annual additive cost. The price
per ton of limestone includes the following:
(1)
(2)
(3)
(4)
FOB limestone cost
Transportation cost
Incremental grinding cost,
Cost of disposal
if any
The capital and operating costs for the disposal system are not
included in the nomograph costs as separate items. These costs were handled
on a cost-per-ton of limestone used. The Kellogg Company uniform estimate
method used a cost of $0. 50/ton for all processes which required the disposal
of bulk waste. In particularly difficult locales which require more costly
methods of disposal, the cost per ton may go as high as $1. 00 to $1. 50/ton.
Since limestone cost calculations are also required for the wet limestone
process, the graph is positioned to be utilized for both limestone processes.
All other operating costs other than the cost of limestone and the
incremental grinding costs are included in a single curve (Graph D). The
major portion of these other costs (which include labor, utilities, plant
-------
IV - 11
overhead, maintenance, etc.) is dependent on the quantity of limestone used.
Thus, the parameters used to scale "other operating" costs are energy
utilization and the L UF.
The total operating cost is found by adding the output of graphs
C, D, E, and F.
2.
Wet Limestone Process
a. Process History. The limestone- injection/wet- scrubbing process
for S02 control was first researched in the United States by Wisconsin Electric
Company and Universal Oil Products Company in 1963 and 1964.
In August 1967, Combustion Engineering received a contract from the
Union Electric Company,on the strength of large-scale tests run at the St.
Clair station of Detroit Edison Company, to install a full- scale combination
wet-dry dolomite process on the 140 Mw unit No.2 at the company's
Meramec Station in St. Louis, Missouri. The unit has been operating only
intermittently because of mechanical problems. Combustion Engineering
installed another system at Kansas Power and Light IS 125 Mw unit No.4
at Lawrence, Kansas. This unit achieved continuous operation in 1970
(Ref. 9).
Both the Meramac and Lawrence systems have recently been
down for modifications based on previous tests. Both are now in operation.
. In addition, Kansas Power and Light is just completing its 420 Mw Lawrence
Unit No.5 which is equipped with the same S02 removal system.
Bechtel Corporation, under contract to TV A and APCO, is cur-
rently installing a major test facility at a 175 Mw unit at Shawnee near
Paducah, Kentucky. This facility, which is also capable of testing dry lime-
stone systems, is being designed to test almost unlimited variations in equip-
ment. Different wet liquor circuits, wet scrubbers, and limestone pre para-
tionsubsystems can be tested in various combinations.
-------
IV - 12
b. Process Description (Refs. 10-13). The Combustion Engineering
process, which employs furnace injection of an alkaline earth followed by wet
scrubbing ,is shown in Figure IV - 2. The pulverized additive calcines in the
furnace and reacts with the combustion gases to form compounds of calcium
and manganese. This removes 10 to 20 percent of the sulfur oxides, including
most of the 803.
Flue gas containing unr-eacted 802 and calcined additive then passes
through one or several air heat exchangers and into the wet scrubber. The
scrubber may be of several designs. The more promising are the venturi,
mobile-packing, and the flooded-bed scrubbers. Because particulate removal
is also a requirement, these scrubber systems are being evaluated not only
for their 802 removal effectiveness, but also for their ability in removing
particulates.
In the scrubber, the calcined additive that had not combined with
802 in the furnace reacts with the water to form a slurry of reactive milk-of-
lime (a suspension of calcium hydroxide in water). The process is
CaO(8) + H20 -> Ca(OH)2(8) + 16 Kcal
It might be noted that this process is accompanied by a threefold expansion in
volume.
The Ca(OH)2 (and Mg(OH)2) reacts with the remaining 802 to form
sulfates and sulfites of calcium and magnesium. Overall 802 removal efficiency
. of about 85 percent is typical. At the same time, water entrainment of fly ash
removes particulate matter. Additive and slurry flow rates control the scrubber
reaction.
The solution containing the reacted materials drains out the bottom
of the scrubber to a tank (clarifier) or pond where the particulates settle. The
clarified slurry is then available for recirculation.
For improved utilization of additive, provisions can be made for
direct recycling of a portion of the scrubber effluent to above-bed sprays. The
cleansed flue gas passes through a mist eliminator for removal of its remaining
water and is then reheated. Reheat is necessary to protect upstream equipment
such as induced draft fans and also to obtain better air dispersion of the effluent.
-------
From Yard Storage
H
o
P
P
E
R
Furnace
To Stack
Stack Gas
Reheat~
Compressor
Limesto e
Coal
Figure IV-2.
Stac k Gas
Reheat
Syste ms
Air
eater
Heat Exchange r /"
- - - -Gas or Oil
J/lli/I/I/If/l_Mist Eliminator
Se ttling Tan~
or Pond
Wet Limestone Injection Process
Recycle
and
Makeup
Water
To Disposal
H
<
I
......
""
-------
IV - 14
Combustion Engineering operations for 1969 at Meramec and
Lawrence, each cleaning 300,000 to 400,000 cfm of gas, revealed information
and pI10blems not evident during pilot plant operation. As a result, these sys-
tems have been modified in the areas of additive injection, gas distribution,
and water control.
Operations when firing 3. 4 percent sulfur coal show an 802 emis-
sion rate equivalent to burning coal with less than 0.5 percent sulfur (.....85 per-
cent efficiency). Greater than 99 percent of the particulate matter can be re-
moved and nitrogen oxide emissions are cut by as much as 30 percent.
c. Process Economics. The wet limestone system requires relatively
few major pieces of equipment. Its capital cost per kilowatt is intermediate
between the dry limestone system and the more sophisticated by-product sys-
terns. The major capital cost element is the flue gas scrubber / demister com-
ponent which amounts to nearly half of the equipment costs. The flue gas flow
rate is the major determinant for sizing the scrubber. The next most costly
items include the disposal/ settling pond and associated piping, pumps, and
controls. The costs of these components are dependent upon the distance the
pond is placed from the power plant, the geology at the pond site, the sulfur
content in the fuel, the stoichiometric ratio, and energy utilization of the plant.
The capital costs of these components are not included in the nomographs since
process wastes were accounted for by charging the process on a cents/ton of
waste basis. Thus, the resultant capital cost estimate is 15 to 30 percent
lower than the estimate which includes settling pond costs. The limestone
storage, handling, and processing equipment make up the last major 8quip-
ment items. These components typically amount to about 15 percent of the
. capital investment.
The nomograph was prepared using four expense categories
including fixed cost, reheat credit, additive cost, and other costs. The fixed
cost is the annual proration for all capital equipment and installation costs
excluding settling pond costs. The additional labor and materials required
for a retrofit installation were not factored into the fixed costs. If an existing
plant was fitted with a process, the fixed cost may increase by 10 to 20 percent
depending upon the modification required to install the process equipment.
-------
IV - 15
The reheat system may either be a gas-gas or gas-liquid-gas heat
exchanger system or it may use a direct natural gas heater. The estimate
used for nomograph preparation includes the capital and operating cost for a
gas-liquid-gas exchanger sized to obtain a 2500F stack temperature downstream
of the process equipment. This is the maximum temperature required to
assure reasonable plume dispersion characteristics. Circumstances
may allow a reheat temperature below 250oF. A credit would have to be
applied to account for the reduction in cost of the reheat system. The size of
the credit depends upon flow rate and reheat temperature. The incremental
fixed cost (assuming 15 percent fixed charge rate) and variable costs were
estimated for various exchanger capacities and reheat temperature resulting
in a credit nomograph which can be utilized if a reheat temperature of less
than 2500F is used.
The additive cost is determined by the same method used in the
dry limestone process. Because the wet limestone method involves higher
fixed charges and has a lower limestone use factor I the significance of the
additive costs are less significant. At the 500 Mw plant size. the additive cost
will typically run about one- third of the total operating cost.
Other operating costs include labor. plant supplies. utilities. and
indirect costs involved in operating the process. As in the dry limestone pro-
cess. the major variables are energy utilization and the limestone use factor.
Other costs typically amount to about one-third of the operating cost of a
500 Mw plant as shown in the example presented previously. Nomograph 3
calculates the above costs. Total operating cost is found by adding the output
of graphs C. D I and F and subtracting the output of graph E.
3.
Mag-nesium Oxide Process
a. Process History. A sulfur dioxide abatement process developed
by the Chemical Construction Corporation (Chemico) utilizes magnesium
oxide in the main reaction. Sulfur dioxide is a by-product which can be con-
verted to sulfuric acid. Extensive pilot plant work has indicated the process to
be technically and economically feasible. A prototype commercial system will
be built at the Mystic No.6 Unit of Boston Edison Company to handle flue gases
from a 155 Mw boiler. Target date for completion and startup is approximately
-------
IV - 16
at the end of 1971, followed by a testing period of 24 months. The nearby
sulfuric acid facility of Essex Chemical Company will be utilized to produce
98 percent concentrated sulfuric acid and to regenerate magnesium oxide.
The magnesium oxide process is also being developed by the
Babcock and Wilcox Company. They built a wet scrubbing pilot plant con-
sisting of three scrubbers. These three scrubbers consisted of a venturi-
type particulate scrubber, a venturi-type absorber, and a tray-type absorber
(floating bed absorber). Many short-term and extended tests were performed
to determine the most satisfactory operating conditions for each scrubber.
b. Process Description. In the process, the flue gas can be first
passed through a venturi-type particulate scrubber to remove greater than
99 percent of the fly ash. At the Mystic site, however, no prescrubber is
employed since the boiler is oil-fired resulting in little fly ash. In Figure
IV - 3, flue gas containing the sulfur dioxide enters a venturi-type absorber in
which it contacts the absorbing liquor in slurry form. The solids in the
absorbing liquor are mainly magnesium sulfite, magnesium sulfate, and
magnesium oxide (insoluble). MgO slurry is added as makeup to the venturi
absorber to maintain the proper reagent concentration. The clean flue gas
leaves the absorber system through a separator and is mixed with hot flue
gas from the dryer and discharged to the atmosphere through a stack.
A small stream of liquor from the absorption system is pumped
to a centrifuge in which MgS03 and MgSO 4 crystals and unreacted MgO are
separated from the mother liquor which is returned to the absorption system.
It is desired to produce a cake with minimum moisture content. The centri-
fuged wet cake is dried in a direct- fired drier to remove the moisture.
Anhydrous dry crystals from the drier are then calcined (decomposed by heat)
in a direct-fired calciner to regenerate magnesium oxide and sulfur dioxide.
The regenerated MgO is returned to the MgO makeup system, to which makeup
water is added to produce makeup slurry for the absorption system. Additional
MgO is also added to the makeup tank to compensate for process losses.
The reactions (Ref. 14) taking place in the major components of the
Chemico process are summarized as follows:
-------
@HEMICO
M 9 0 SCRUBBER SYSTEM - 0 I L F I RED BO I LERS
(Ref. Chemical Construction Corporation)
ve nturi ~
MgO Scrubber Absorber
Elevator FD Fan
Air
0 Fuel
MgO Dryer
To Centrol
Storage Process Plan t
MgO from Central
Process Plant
MgO
Hopper
H20
Recyc Ie
. MgO
Make-up Pump
Dust
Collector
Stack
Mother
Liquor
Tank
~
Centri fuge
Recyc I e
Pump
Mother
Liquor Pump
Figure IV - 3
H
<:
I
.....
-J
-------
IV - 18
Main Reaction:
MgO + S02 + 6H20
Side Reactions:
MgS03 + S02 + H20
Mg(HSOS)2 + MgO
MgO + S03 + 7H20
2MgSOS + 02 + 14 H20
Drier:
MgSOs .
MgS04 .
Heat>
Heat>
6H20
7H20
H20
Heat>
Calciner:
MgS03
2MgS04 + C
Heat>
Heat>
>
MgSOs . 6H20
::>
lVlg(HSO 3) 2
2MgS03 + H20
MgS04 . 7H20
2(MgS04 . 7H20)
>
>
>
MgS03 + 6H20 t
MgS04 + 7H20 ;\.
H20 t
MgO + S02'"
2MgO + 2S02':'
+ C021
A schematic of the magnesia process proposed by Babcock and
Wilcox is shown in Figure IV -4, (Ref. 15). This co nfiguration is assumed to
be integrated into an 800 Mw boiler system.
The system starts at the plenum which collects the flue gas leaving
the primary and secondary air heaters. The flue gas passes through the fol-
lowing equipment which is arranged in parallel: four venturi particulate
scrubbers, four sulfur dioxide absorbers, four induced draft fans, and two
steam coil gas repeaters. The gas is then discharged through a stack. The
entrances to the stack are considered terminal points.
The gas entering the venturis is mixed with water which contacts
and removes the entrained fly ash. The resulting fly ash slurry is trans-
ferred to a thickener for separation. The clarified overflow is returned to
the venturis for reuse while the underflow is pumped to filters. Here, the
fly ash is washed with fresh water to minimize the amount of soluble sulfur
and magnesium compounds that leave the system with the fly ash. Pumps
-------
Boiler
Air
Heater
Coal
Coal
Fly Ash
Scrubber
Air
Forced
Draft Fan
Fly Ash
Dewatering System
To Ash Disposal
78% H2S04
By- product
S02
H2S04
Plant
Figure IV-4.
S02 Absorber
Flue Gas
Reheater
Induced
Draft Fan
Spent Scrubbing
Slurry Dewatering
System
Spent Scrubbing
Slurry Drying
System
Magne sia
Regeneration System
Babcock & Wilcox MgO Process
Stack
1-4
~
,
......
CD
-------
IV - 20
have been included to sluice the fly ash to a pond.
pumps is the terminal point of the ash system.
The particulate-free flue gas passes to the absorbers where sulfur
dioxide is removed by reacting with magnesium sulfite. The resulting mag-
nesium-sulfur compounds are dewatered by thickening and centrifuging. The
liquor which is separated from the slurry in the thickener and centrifuges is
returned to the absorbers. Magnesium oxide. which results from the regener-
ation of the magnesium-sulfur compounds. is added to the absorbers to replace
the magnesium which reacts with sulfur dioxide and is removed to the thickener
as a sulfide. The solid materials leaving the centrifuge are dried and then
decomposed to yield magnesium oxide and a gas stream containing sulfur
dioxide. The product gas leaving the reactor is cooled. filtered. and leaves
the system. It is then converted into sulfuric acid at an acid plant.
The discharge of these
In the pilot plant studies performed by Babcock and Wilcox (Ref. 16)
the efficiencies of the venturi-type and tray-type absorbers were measured.
These absorber types are employed in the Chemico system and the Babcock
and Wilcox system. respectively. The 802 removal efficiency for the venturi
absorber is presented in Figure IV -5 as a function primarily of venturi pressure
drop. It was reported that. except for pressure drop. the 802 removal per-
formance of the venturi absorber showed a minimal dependence upon the
parameters examined. The other variables considered were liquid-to-gas
ratio. MgO/802 stoichiometric ratio. pH of the scrubbing slurry. the presence
of fly ash. the presence of moderate concentrations of Mg804' slaking or con-
version of MgO to Mg(OH)2 before injection into the venturi. and residence
time of the liquid. The correlation between efficiency and pressure drop in
the figure was judged as being based on insufficient data since there were only
three data points at elevated pressure drop.
. The 802 removal efficiency for the tray-type absorber or the
floating bed absorber is shown in Figure IV - 6. The 802 absorption data
correlated well in terms of two operating parameters. L/G and spray slurry
composition. The slurry composition. which is dependent upon the ratio of
MgO/802 and solubility of Mg803' can be expressed indirectly in terms of
the bisulfite concentration. Mg(H803)2 or slurry pH. Thus. the figure illus-
trates the absorption efficiency as a function of L/G and bisulfite concentration.
-------
-
~
-
:>,
C)
!:i 90
Q)
....
C)
....
.....
.....
fi1
!:i
0
....
+-' 85
0-
s....
0
U)
..0
~-
C'J 80
o
r.n
Figure IV- 5.
IV - 21
95
J
J
,
~ 'f .!J I 1[ ,
~r-pH > 7. O. J
II
~
I
~
,.
~
I
f-- ~ '-I- ~ -
-- -
--l.. --
75
70
o
2
4
6
8
10
Venturi Pressure Drop (inches, w. g. )
802 Removal Efficiency for Venturi-Type Absorber
-------
, ,!
80 :
I
i
70 '
,
60 :
,
40 I
20
0
0
IV - 22
!)!).6
99.4
99.2
D!)
!)B
....,
!::
-,
C)
!::
o
e
-------
IV - 23
c. Process Economics. The variations in capital and operating costs
for the magnesium oxide are illustrated on nomograph pages 14-16. Each of the
figures included in the nomograph is discussed in the succeeding paragraphs.
The five fixed capital cost estimates performed by the Kellogg
Company apply to a Babcock and Wilcox system. Major equipment costs for
this system are contributed by the venturi particulate scrubbers, tray-type
absorbers, centrifuges, driers, gas reheaters, calciners, and the H2S04
plant. An estimate was also performed by Hittman Associates at a plant
capacity of 2000 Mwe and a 3. 5 weight-percent sulfur coal. The resulting
fixed capital cost of $42. 8 million was an extension of the base case computed
by Kellogg. One additional cost coordinate occurs for the trivial case at zero
fixed capital cost and flow rate. Additional constant sulfur curves were obtained
by interpolation and extrapolation of the data so that the range of sulfur content
in fuels from o. 5 through 5. 0 lb /106 Btu was covered. The final curves which
were plotted on log-log graph paper allows one to obtain a fixed capital cost
at a given flow rate and fuel sulfur content. The fixed cost in millions of
dollars per year is then obtained for the fixed capital cost and a predetermined
fixed charge rate.
Raw material costs are due to the annual requirement for makeup
magnesium oxide. Makeup MgO quantity depends upon losses taking place in
the drier and calciner or regeneration operations of the system. Magnesia
requirement also varies with plant capacity and energy utilization. Magnesium
. oxide quantity has been assumed to vary proportionally with energy utilization
and is based on the annual requirements presented in the base case.
The by-product credit for 98.5 percent concentrated H2S04 is
shown to be a linear function of sulfur removed. The sulfur removed by the
process is the product of three quantities: 802 removal efficiency, energy
utilization, and sulfur content. Appendix Table B-9 indicates that the assump-
tion throughout of a constant 802 removal efficiency of 95 percent, exhibited
by the three 3.5 percent sulfur cases, will introduce little error in the sulfur
removed quantity. The quantity of 98.5 percent H2804 is then a constant multiple
of the sulfur removed quantity.
Other operating costs include utilities, direct labor, supervision,
maintenance, and overheads. These are presented as a function of energy
utilization and the sulfur content of the fuel ranging from 0.5 through 5 lb/106 Btu.
-------
IV - 24
4.
Catalytic Oxidation Process
a. Process History. The catalytic oxidation process ("Cat-Ox")
developed by Monsanto for the removal of sulfur oxides from stack gases is
essentially an adaptation of the well-known contact process for the manufacture
of sulfuric acid. The gas containing both sulfur dioxide and oxygen is passed
through a fixed bed of catalyst at an appropriate temperature and most of the
S02 is oxidized to S03. The gas is then passed through an absorption tower
where the S03 is absorbed in recirculated sulfuric acid. The by-product for
the Cat-Ox process is 78 percent concentrated sulfuric acid.
For power plant applications, the Cat-Ox process is offered in two
configurations, one for use in new plants and one for previously existing plants.
Several years of operating experience were gained on a now-dismantled pilot
plant at Metropolitan Edison's Portland, Pennsylvania, station. A retrofitted
system is planned for Illinois Power's 100 Mw Wood River 4 unit. Initial
operation is expected in mid-1972.
b. Process Description. For the configuration that is integral with
a new plant (Figure IV -7), components that are required in common for both
the power plant and the Cat-Ox process are the electrostatic precipitator, fly
ash handling system, economizer, air heater, stearn lair heater, induced
draft fan, and the forced draft fan. The electrostatic precipitator operates at
a high temperature and is required to be highly efficient. In the converter, a
single pass through the vanadium pentoxide catalyst bed results in 90 per cent
conversion of S02 to S03. The air heater recovers less of the heat from the
flue gas temperature than is done in a conventional power plant. The outlet
flue gas temperature at the air heater is determined by the necessity for re-
maining above the acid dew point to avoid corrosion. Combustion air is pre-
heated first in the fluid/air heater and second in the stearn/air heater. The
fluid/ air heater is employed to recover additional heat from the recirculated
acid which is cooled after each pass through the absorbing tower. The absorp-
tion of the sulfur trioxide in the cooled, recirculated acid is accomplished in a
packed tower. The sulfuric acid mist formed during the cooling of the gas,
plus any entrained droplets of the circulating acid carried out of the tower,
are recovered in a fiber-bed mist eliminator. The minimum outlet gas tem-
perature at the absorbing tower is set by the concentration of the sulfuric acid
produced (78 percent).
-------
~r---: .
, I
, ,
Flue Gar;; ". : I
~~~l~rs;p . .t t I Ic t I. L .1 i Air .1 InDduCfetd i Absorbing, 11 Mist 1~~U~t:::
recipi a 0 - onver er ---Dconomizer"""'--'! H t. ra --! T ~ El. . t i .~
, 8g0 18goi .. ~501 ea er! Fan.. t40i ower:: ImIna oJ 2380F
Ash DIsposal Area . 4300F . - - _. - .
Boiler Feedwater :
---- 1--. ---1! I
St / A. ! i ACId i ! Process I
~~~ter Ir L;,. ICirculating, ~: Acid'
, Pump; i Tank
~ ._-~ '
FIUid/A~ll. ~!J AC~ r prOduc~OOling Water
~ cir~~~~tionh Circulation: Acid
Heater. I I '
I Cooler Cooler Cooler
Boiler Feedwater
~ ---------.--...- ..__.
.. - . .-. u.---.... ,
i
Condensate
-------.
Air for Insulators
,~
Combustion Air
Steam
LEGEND:
Primary Flow
-- - Intermittent Operation
,
1-
Air to Fan
l
CondensatelJ
Heat I
. Exchanger I
Condensate
-_._,..".~---_.. ----
Figure IV -7. Integrated Catalytic Oxidation System
For New Power Plant
Cooling Water
H
<:
I
t-.:)
OJ
-------
IV -26
The Cat-Ox reheat system for an existing power plant is shown
schematically in Figure IV - 8. It is additive instead of integral. The reheat
system could be bypassed entirely without affecting the operation of the power
plant except d ur ing a relatively short pe r iod in which the final tie - ins are made.
This period could coincide with the annual shutdown of the power plant for
overhaul.
The reheat system takes the gas from the discharge of the existing
induced draft fan. A new precipitator is added to reduce the quantity of dust
to a tolerable level, which is the same as in the integrated Cat-Ox system. A
. .
new induced draft fan is added to move the gas through the Cat-Ox system.
The relatively cool (3250F) flue gas must be heated to a temperature high
enough to permit catalytic oxidation of the sulfur dioxide. Part of this pre-
heating is accomplished by exchange of heat from the converter exit gas in the
gas heat exchanger. The rest of the preheating is accomplished by direct
injection of hot combustion gases from the reheat furnace, which can be fired
with No.2 fuel oil of low ash content. The hot gas leaving the converter
passes through the gas heat exchanger where it heats the incoming cool feed
gas and is itself cooled to 4500F before it enters the absorbing tower. From
this point on, the arrangement and operation are essentially the same as in
the integrated Cat-Ox system. The temperature of the gas leaving the heat
exchanger, 450oF, is set as in the integrated system by the necessity of
staying above the sulfuric acid dew point. In the gas heat exchanger, as in
the air heater of the integrated system, it is necessary to avoid cold-end
corrosion. The temperature of the incoming cool flue gas must therefore be
raised somewhat and this is accomplished by the recycle of a portion of the
hot exit gas to the inlet of the exchanger. The sizes of the new induced draft
fan and the gas heat exchanger must be increased accordingly to handle the
greater total flow of gas. The acid circulated to the absorbing tower is
cooled directly with cooling water and there is no recovery of the low-level
heat. The product acid is of 78 percent concentration, as in the integrated
system.
For a new or integrated system, 802 removal efficiency depends
upon the converter temperature and the contact time in the catalyst bed.
Reference 17 indicates that this efficiency is 90 percent at the nominal con-
verter temperature of 850oF. For a retrofitted process or reheat system,
-------
From
I. D.
Fan
Air for
Insulators
32
o
Precipitator 325
of
LEGEND:
-----
New
Induced
Draft Fan
Primary Flow
Intermittent Operation
3£0
of
Gas Heat
Exchanger
Absorbing
Tower
Reheat
Furnace
Combustion
Air Fan
Fuel Oil
Storage
1
,.--
I
I
I
I
I
I
I
Process
Acid Tank
Converter
8700F
Acid
Circulation
. Coolers
Mist
Eliminator
- - - - - - - - - - - - - - - -- - _.J
..
Product
Acid
Cooler.
Figure IV- 8. Catalytic Oxidation Reheat System for Existing Power Plant
7800F
4500F
Effluent
Cooling Water
Stack
Cooling Water
Product
Acid
Storage
Flue Gas to
Atmosphe re
2360F
Product
Acid to
Loading
......
<:
I
I:-.:J
-J
-------
IV - 28
an overall S02 removal efficiency of 90 percent is specified by Monsanto.
Graphical presentations of the S02 removal efficiency are unavailable at this
time for the Cat-Ox process.
c. Process Economics. The variations in capital and operating costs
for the Cat-Ox process are illustrated by a series of figures on nomograph pages
17-19. The fixed capital cost estimates generated by the Kellogg Company
apply to a new process or an integrated Cat-Ox system. Additional estimates
were computed by Hittman Associates at plant capacities of 1500 and 2000 Mw
and a 3.5 weight-percent sulfur coal. The resulting magnitudes were $57.2
million and $69.6 million, respectively. These were extensions of the base
case. The family of fixed capital cost curves, spanning the range of sulfur
content from 0.5 through 5.0 Ib/106 Btu, is closer together than for the MgO
process. The capital costs for the Cat-Ox process are thus less sensitive to
variations in sulfur content found in fuels. A fixed cost in millions of dollars
per year is obtained for a given fixed capital cost and a given fixed charge
rate.
Raw material costs are due to the annual requirement for vanadium
pentoxide (V 205)' V 205 replacement cost is incurred by a loss of the catalyst.
This is caused by the cleaning operations necessitated by the gradual buildup of
fly ash in the catalyst bed. Even though an electrostatic precipitator with 99.6
percent efficiency is used, enough fly ash enters the bed to require four cleanings
per year. A 2.5 percent loss of catalyst occurs during each cleaning. The
replacement quantity is dependent upon the energy utilization and the fly ash
content of the fuel. Since the exact functional dependency is unknown, the base
case operating cost estimate was used to establish a family of 450 lines for the
annual volume of V 205' The ash content of the coal on the as-received basis
was 14.5 percent.
The by-product credit for 78 percent concentrated H2S04 is shown as a
function of sulfur removed. The sulfur removed by the Cat-Ox process is
presented versus energy utilization and sulfur content after assuming an S02
removal efficiency of 90 percent. Table B-10 of Appendix B indicates that
this assumption will introduce a maximum error of under 10 percent in the
sulfur removed quantity. The quantity of 78 percent H2S04 is then a constant
times the sulfur removed quantity.
-------
IV - 29
Other operating costs include utilities, direct labor, supervision,
maintenance, and overheads. A credit for stack gas cooling has been accounted
for in the process. Other operating cost exhibits negligible variation with
changing sulfur content. Thus, a single line results after plotting other
operating costs versus energy utilization.
5.
Modified Chamber Process
a. Process History. The catalytic chamber process enables the
removal of both the oxides of sulfur and nitrogen from power plant flue gas.
The process is a modified version of an earlier chamber method to produce
sulfuric acid. Commercial grades of sulfuric and nitric acids are by-products
of the process. The development of the process was initiated in October 1967
and proceeded until October 1970 under two contracts for the Environmental
Protection Agency. The effort by Tyco Laboratories, Inc., is continuing
under a new contract to the Air Quality Office. To date, the experimental
work has consisted of a laboratory scale investigation of the separate process
stages and a process evaluation on a continuous basis in a 10 scfm pilot plant.
b. Process Description. A schematic diagram of the Tyco catalytic
chamber process (Ref. 18) is shown in Figure IV - 9. The flue gas, fed to the
process, is mixed with recycled nitrogen dioxide which oxidizes the sulfur
dioxide to sulfur trioxide in a reactor. Water vapor in the flue gas stream
converts the 803 to sulfuric acid. The reacted gases pass through the scrubber
which removes the newly formed acids and the oxides of nitrogen while venting
the cleaned, wet stack gas to the atmosphere. Fly ash scrubbed from the flue
gas is filtered from the acid stream. The nitrogen oxide- bearing sulfuric acid
is then passed onto the catalytic stripper where it is contacted with air in the
presence of activated carbon. Here the nitrosylsulfuric acid is catalytically
oxidized to H2S04 and N02. The N02 - bearing gas stream is fed to a nitric
acid production stage where the excess N02 is recovered. The H2S04 leaving
the stripper is recycled to the scrubber after the by-product acid is removed.
This by-product acid is equivalent to the amount of 802 recovered from the
flue gas in the scrubber after oxidation by N02.
-------
Gas
NO
N02
S02
Gas to stack
2500F
7% H20
150 ppm NO
x
150 ppm S02
Reactor
HN03 Absorber
Air
NO
Flue gas
3000F
O. 3% S02
O. 06% NO
x
By- Product
60% HN03
Gas
N20S
H2 SO 4
Air N02
H20
Activated
Charcoal
Heated
Air or
Flue Gas
By- Product
80% H2S04
80% H2S04
High
Temperature
Scrubber
80% H2S04
HNS05
Catalytic
Stripper
1-1
<
I
c..:I
o
Figure IV- 9. Tyco Catalytic Chamber Process Flow Diagram
-------
IV - 31
The reactions taking place in the major components of the pro-
cess are summarized as follows:
Reactor:
S02 + N02 -> S03 + NO
<-
S03 + H20 -> H2S04
<-
2NO + 02 -> 2N02
<:
NO + N02 -> N203
<-
Scrubber:
N203 + 2H2S04
Catalytic Stripper:
->
<-
2HNS05 + H20
4HNS05 + 02 + 2H20
->
<-
4N02 + 4H2S04
Nitric Acid Production:
3N02 + H20
->
2HN03 + NO
The S02 removal efficiency for this process is a function of
reactor residence time, reactor temperature profile, and the N02:S02 ratio
into the reactor. From Reference 18, following experimentation with the S02
reactor, a residence time of approximately 15 seconds was required to yield
an S02 removal efficiency of 95 percent from flue gas containing O. 3 percent
S02 by volume. The N02:S02 ratio into the reactor is nominally 2.0, but it
can range from less than 2.0 up to 2.2. The reactor temperature profile used
in the pilot plant was relatively flat since the temperature difference between
inlet and exit was small, 3350F and 372oF, respectively (Ref. 19). Thus, the
reactor operating temperature used in obtaining the data was approximately
constant. Following further operation of the 10 scfm miniplant, a better
definition should result for the parametric variation of S02 removal efficiency.
c. Process Economics. The variations in capital and operating costs
for the modified chamber process are illustrated by a series of figures on nomo-
graph pages 20-22. The fixed capital cost estimates generated by the Kellogg
Company apply to a new Tyco process employing a packed scrubber as opposed
to a plate tower scrubber. Major equipment costs for this system are contri-
buted by the flue gas blowers, the reactors, the packed scrubbers, and the
-------
IV - 32
catalytic strippers. An additional capital cost estimate was computed by
Hittman Associates as being $48. 3 million at a plant capacity of 2000 Mw and
a 3.5 weight-percent sulfur coal.
Fixed capital cost versus flue gas flow rate is represented by a
single line. The costs exhibited negligible variations resulting from pertur-
bations in sulfur content. The annual fixed cost in millions of dollars per
year depends upon the capital cost from the first portion of the graph and the
fixed charge rate.
Raw material costs are due to the replacement costs for the acti-
vated charcoal bed in the catalytic stripper and the limestone required to
neutralize the filter cake. The parametric study of Reference 19 showed that
the annual quantity of granular charcoal catalyst varied in direct proportion to
energy utilization. It was not indicated as being a function of the fuel sulfur
content. Therefore. the annual cost to replace the charcoal beds is shown
parametrically to account for variations in unit cost alone. The annual cost
of limestone is small (approximately one percent of direct costs) and conse-
quently is lumped into "other operating costs. "
Credits are obtained for the by-product acids. 80 percent concen-
trated H2S04 and 60 percent concentrated HN03' It was shown in Reference 19
that the annual quantity of sulfuric acid was a function bf energy utilization and
sulfur content of the fuel. The quantity is shown as varying proportionately to
increases in energy utilization and variations in sulfur content. Therefore.
the annual revenue for by-product H2S04 is obtained for a given energy utili-
zation. sulfur content. and net sales revenue. The sulfur removed at an S02
removal efficiency of 87.5 percent can also be obtained from the graph.
The annual quantity of by-product nitric acid is a function of
energy utilizatio n and the NOx content of the flue gas. The HN03 quantity is
proportional to energy utilization. This quantity was shown (Ref. 19) not to
be directly proportional to the NO content of the flue gas. The concentration
x
of nitrogen oxides in the flue gas is a function of the boiler design and operating
temperature. excess air. fuel type. and firing methods. For coal-fired units.
a representative range of NO content in the flue gas by volume would be 400
x
to 1200 ppm. The Kellogg base case assumed a value of 500 ppm. This
appears to be reasonable when compared to the magnitudes presented in
Reference 20.
-------
IV - 33
Other operating costs include utilities, direct labor, supervision,
maintenance, overheads, and limestone material. A credit for boiler feed-
water heating has also been accounted for. Other operating cost exhibits insig-
nificant variation with changing sulfur content. Thus, a single line results
after plotting it versus energy utilization.
6.
Molten Carbonate Process
a.
Process History.
The molten carbonate process to remove sulfur
oxides from power plant flue gases is being developed by the Atomics Inter-
national (AI) Division of North American Rockwell. The effort was initiated
on June 1, 1967, under contract with the National Air Pollution Control Admini-
stration of the Department of Health, Education, and Welfare. Bench studies
and conceptual testing for the molten carbonate process have been performed.
However, construction of a pilot plant for the process is only in the planning
stages. During March 1969, an advisory committee of the National Academy
of Engineering recommended that the pilot plant be built to treat a side stream
of gases from an existing coal-burning boiler. It is anticipated that a site may
soon be selected.
b. Process Description. In the process shown in Figure IV-10, flue
gas containing sulfur oxides is scrubbed in absorbers with a molten mixture
of lithium, potassium, and sodium carbonates (M2C03)' The process tempera-
ture is nominally 850oF. The melt, containing sulfite salts, is regenerated
. by heating to 15000F where the sulfates, formed by disproportionation of the
sulfites, can be reduced to sulfides with coke. The sulfides in the melt are
converted back to carbonates, with the formation of hydrogen sulfide, by
reaction at 850-900oF with carbon dioxide and steam. The H2S is sent to a
Claus Plant for recovery of by-product sulfur. A description follows of the
reactions taking place in the major components of the process.
Flue gases tapped before the economizer section of the boiler are
passed through a bank of high temperature electrostatic precipitator sections
operating in parallel. Precipitator efficiency is 99.5 percent. The gases are
sent from each precipitator to a corresponding absorber for removal of the
sulfur oxides by the molten carbonate stream. The return flue gas to the
economizer section of the boiler has had greater than 90 percent of the original
SO removed.
x
-------
M2C03
Makeup
Pump'
Return to
Boile r
t
Flue Gas
850°F
Absorber
850°F
Recycle
M2C03
M2S03
M2S04
Flue Gasl
With S02!
and I
Fly Ash L'
Hot Fly Ash
Precipitator
Pump
Flue Gas
(From Boiler)
-850°F
H2S to Claus Plant
-8500F I 2
M2C03
M2S
-850°F
Filter J
850°F
=r=---
Heat I
Exchanger!
} Return to Process
......
<
I
(A
~
C02
(
Pump'
I
~Air
Coke Filter
Cake, Coke
Residue, and
M2C03
M2S
Reducer
I
-1500oF . ~ Coke
M2C03
M2S03
M2S04
Filter
",850°F
Fly Ash
Filter Cake
Fly Ash and
M2C03 }
M2S03
M2S04
Return to Process
Figure IV -10. Molten Carbonate Process Flow Diagram
-------
IV - 35
Absorber reactions are:
S02 + M2C03
S03 + M2C03
->
M2S03 + C02
M2S04 + C02
->
where M represents the eutectic mixture of sodium, potassium, and lithium
cations. Some of the M2S03 may be oxidized by oxygen in the flue gas as
follows:
2M2S03 + 02-> 2M2S04
All sulfides are also oxidized as follows:
M2S + 202
-> M2S04
Fly ash is removed from the process by two filter units in parallel.
One unit is in operation while the other unit is being drained and a "dry" ash
cake is discharged. The fly ash filter cake is composed of a mixture of molten
salts and solids consisting of KCI, fly ash, and a trace of coke. The fly ash
filter cake is sent to the Li2C03 recovery process which recovers over 80
percent of the Li2C03. At every discharge, a small amount of reactant is
lost requiring a carbonate make-up.
Each reducer consists of oxidation and reduction zones with liquid
passages between the two compartments for internal melt recirculation. The
endothermic heat for reduction and the sensible heat required to raise the
melt, coke, and air to the reduction temperature of 15000F are provided by
the oxidation of M2S in the oxidation zone of the reducers. At 1500oF, the
M2S03 in the melt disproportionates into M2S04 and M2S. The M2S04 is
reduced in the reduction zone to M2S by the carbon contained in the coke.
Part of the reduction zone melt is recycled to the oxidation zone to oxidize
the M2S and provide the necessary heat. The oxidation zone melt also recycles
back to the reduction zone to convert the M2S04 into M2S. Reducer reactions
are as follows:
Oxidation:
M2S + 202
C + 02
->
M 2804
C02
->
-------
IV - 36
Reduction:
M2S04 + 2C ->
Disproportionation:
M2S + 2C02
4M2S03
Coke Sulfur Recovery:
->
3M2S04 + M2S
2S + C + 2M2C03 -> 2M2S + 3C02
The melt from the reducer at 15000F is cooled.
Melt from the reduction step flows to one of two parallel coke
filters where the unreacted coke is removed from the process melt. The
filtered melt is sent to the regenerator. The coke filter cake is composed
of coke plus traces of KCI and fly ash. The coke filter cake is sent to the
Li2C03 recovery process. Additional carbonate make-up is needed to replace
the nonrecovered carbonates in the coke filter cake.
The regenerator isa tower consisting of bubble caP trays. The
melt is added to the top of the column and is contacted with the C02 and H20
rising up the tower. The regeneration is exothermic and coolers are provided
to remove this heat. The regeneration reaction is:
M2S + C02 + H20
->
M2C03 + H2S
The gases from the regenerator containing H20, C02' and H2S are sent to a
Claus Plant for recovery 'of by-product sulfur.
The absorber design criteria proposed by AI is incorporated into
Figure IV -11 as presented by Singmaster and Breyer (Ref. 21). In addition,
a maximum gas velocity of 25 fps is proposed. In the figure, the cross-
hatched area represents the limiting region for operation. On the abscissa is
the inlet molar ratio defined to be the total mols of molten carbonate per hour
to the absorber divided by the mols of SO per hour entering in the flue gas.
x
On the ordinate is the molten carbonate concentration at the absorber outlet.
Recycle is defined on the figure. The desired target operating condition has
the coordinates corresponding to a recycle of 1:1 and 68 percent carbonate
melt concentration at the absorber outlet.
-------
100
90
L=2724
1 = 2 724
L/G=0.0116
No Recycle
~ 80l
~
~ I
I
.J
Q) 70 L 68%
~ ,------ -~~~22
.
-------
IV - 38
The bounds for the inlet molar ratio are approximately 3.5:1 and
10:1 from the figure. In Reference 22, the 802 removal efficiency was pre-
sented versus the inverse of the inlet molar ratio. It is shown here in Figure
IV -12. These data resulted from experiments employing an absorption column
apparatus wherein operation occurred at low and high (fewer cases) gas velocities
The efficiencies corresponding to these bounds were approximately 93 percent
and 98 percent, respectively. Therefore, the 802 removal efficiency is very
high and is essentially invariant over the range of operating conditions for the
process absorber.
c. Process Economics. The variation in capital and operating costs
for the molten carbonate process are illustrated by the series of figures on nomo-
graph pages 23-26. The fixed capital cost estimates generated by the Kellogg
Company applied to a new process. Additional estimates were computed by
HittmanAssociates at 1500 and 2000Mw. Major equipment costs for this
system are contributed by the high efficiency (Tl = 99.5 percent), high tempera-
ture electrostatic precipitator, the absorbers, the reducers, and the Claus
sulfur plant. A family of curves was obtained for the range of sulfur content
from 0.5 through 5.0 lb/106 Btu. A fixed cost is obtained for a given fixed
capital cost and fixed charge rate.
The by-product credit for sulfur is determined after multiplying
the unit cost of sulfur by the annual quantity of by-product sulfur. The by-
product sulfur is known to be approximately equal to the quantity of sulfur
removed from the flue gas. The quantity of sulfur removed was determined
versus energy utilization and sulfur content at an 802 removal efficiency of
95 percent.
Raw material costs are due to coke and carbonate makeup require-
ments. The annual quantity of coke required as input to the reducer is directly
proportional to the sulfur removed. The annual coke cost is read directly for a
given sulfur removed and coke unit cost. If the unit cost of coke becomes
excessive, alternate low cost raw materials may suffice as sources of the
carbon required for reduction.
The annual replacement cost for makeup carbonate is determined
from nomograph pages 25 and 26 by summing the following three contributions,
respectively:
-------
~ 95-
(I)
t.;
~
rn
UJ
(Ij
o
(I)
..c::
+->
s
o
t.;
.....
'"0 i
~ I
s 90L
(1)
p::;
C'\I
o
'U)
~
s::
(I)
C)
t.;
(1)
P-t
lOO~
. "'"
"",-,
"~
~
'"
O'~~-l
,,",,
U
"-
,:J "'"
I I
0.20 0.30
802/M2C03 Ratio
Figure IV-12. Effect of 802-M2C03 Ratio Upon the
Percent Removal of 802
\..;
.-,
J
85 -
I
o
I
O. '10
I
0.40
IV - 39
----l
I
I
I
!
,
',./
Least squares lines
~""
"
"-
"
. "\....,
-"'"
"
~,
"~
~
'"
J
0.50
0.60
-------
IV-40
(1)
(2)
Makeup cost forM2C03 loss due to fly ash filter cake
Makeup cost for M2C03 loss due to coke filter cake
(3)
Mq.keup cost for M2C03 loss due to KCl filter cake and
KCl production
These costs depend upon energy utilization and sulfur removed (for the coke
filter case case only) shown on the abscissa.
In Reference 21, Singmaster and Breyer developed the carbonate
makeup costs shown in Table IV - 3 for their base case, 800 Mw capacity. Thus,
for a given energy utilization, it can be shown that the makeup cost for loss
due to the fly ash filter cake is:
Cost contribution = (Li2C03 unit cost x Ib of Li2C03 makeup per Ib of
(1)
fly ash + K2C03 unit cost x Ib of K2C03 makeup
per Ib of fly ash + Na2C03 unit cost x Lb of Na2C03
makeup per Ib of fly ash)x Ib/hr of fly ash removed
at fly ash filter x 8760 hr /yr x plant load factor
The Kellogg base case operq.ting cost estimate of Table B-6
e=mployed a fly ash removal rate corresponding to 14.5 percent ash content in
the coal. The removal rate was increased and decreased proportionately in
order to obtain makeup costs corresponding to the 5, 10, and 15 percent ash
contents, respectively. Thus, makeup costs due to losses in the fly ash filter
cake are presented as a function of energy utilization and the percent ash con-
tent in the coaL.
Similarly, for a given energy utilization, the carbonate makeup
cost for loss due to the coke filter cake is:
Cost contribution =
(2)
(Li2C03 unit cost x Ib of Li2C03 makeup per Ib of
unreacted coke + K2C03 unit cost x lb of K2C03
makeup per Ib of unreacted coke + Na2C03 unit
cost x Ib of Na2C03 makeup per lb of unreacted
coke') x Ib/hr of coke removed at the coke filter
x 8760 hr /yr x plant load factor
-------
TABLE IV-3. CARBONATE MAKEUP REQUIREMENTS
ASSUMING 88 PERCENT LI2C03 RECOVERY
(Ref. Singmaster and Breyer Base Case)
Cost Per Pound
Makeup Ratio Required for
Non-Recovered Salt, #/# of
Fly Ash and KC1 Filter
Cake
Coke Filter Cake
KCl Produced
Annual Replacement
for Loss From
Fly Ash Filter Cake
KC 1 Filter Cake
Coke Filter Cake
KCl Production
Total
Annual Cost for Makeup
Due to Losses in
Fly Ash Filter Cake
KCl Filter Cake
Coke Filter Cake
KCl Production
Total
Li2 CO 3
Basis:
Filter Cake Constituent:
Fly ash- 225 lblhr
KCl -403 lb/hr
Removed at fly ash filter
Coke - 800 Ib /hr
Removed at coke filter
K2C03
Na2C03
$0.42
$0.096
$0.028
0.067
0.611
O. 573
0.043
0.368
0.926
0.361
93,000# 843,000#. 791,000#
167,000 1, 510,000 1,417,000
211,000 1, 805, 000 1, 771,000
2,288,000
471,000# 6,446,000# 3,979,000#
TO TA L
$ 39,000 $ 81,000 $ 22,000 $142,000
70,000 145,000 40,000 255,000
89,000 173,000 50, 000 312,000
220,000 220,000 I-<
<:
I
$198,000 $619,000 $112,000 $929,000 ~
......
-------
IV - 42
The unreacted coke Or coke residue removal rate was determined
parametrically in Reference 21 as being directly proportional to plant capacity
and percent sulfur in the coal. The quantity in Table IV -3 is based on the
assumption of complete oxidation of carbon in the coke to provide the reducer
sensible heat and heat of reduction. A conversion of 95 percent of the M2S04
to M2S by carbon is also assumed. Therefore, makeup costs due to losses in
the coke filter cake are finally presented as a function of the sulfur removed
at the removal efficiency of 95 percent.
At constant energy utilization, the carbonate makeup cost for loss
due to the KCl filter cake and the KCl production is as follows:
Cost contribution = (Li2C03 unit cost x lb of Li2C03 makeup per lb of
( 3)
KCl filter cake + K2C03 unit cost x lb of K2C03
makeup per lb of KCl filter cake + Na2C03 unit
cost x lb of Na2C03 makeup per lb of KCI filter
cake + K2C03 unit cost x lb of K2C03 reacted per
Ib of KCl produced) x lb/hr of KCl removed at the
fly ash filter x 8760 hr /yr x plant load factor
The potassium carbonate reacts with chlorides in the flue gas, such as HCl,
which are present due to the chlorine content of the fuel:
K2C03 + chlorides
138. 2 lb /lb mol
>
2KCl +. . .
149. 1 lb/lb mol
Thus, the lb of K2C03 reacted per lb of KCI produced is equal to 138.2/149.1
or 0.926, and this additional makeup ratio is shown in Table IV-3 along with
the others.
The Kellogg base case operating cost estimate of Table B - 6 would
have employed a KCI removal rate corresponding to 0.03 percent chlorine con-
tent in the coal. This rate was gained proportionately in order to obtain make-
up co sts corresponding to the percent chlorine content in the coal of O. 01 percent,
0.05 percent, and O. 10 percent. According to Reference 23, chloride content of
American coals can range up to 0.5 percent by weight. Thus, makeup costs
due to the losses from KCl filter cake and production are presented versus
energy utilization and percent chlorine content in the fuel.
-------
IV -43
The makeup ratios involving Li2C03 incorporate in their magni-
tudes the basic assumption of 88 percent Li2C03 recovery.
The unit costs employed in the above equations for the lithium,
potassium, and sodium carbonates are from Table B-6. These unit costs
were compared to those employed in Reference 21 and those quoted recently
from two local suppliers. Deviations in unit cost were small for Li2C03 and
K2C03. The Kellogg unit cost for K2C03 was approximately one-half of the
three others after comparison. Deviations in unit cost were greatest for the
Na2C03. However, the Na2C03 makeup cost is the least contributor to the
total makeup cost. Therefore, usage of the Kellogg unit costs gives a rea-
sonable representation of the carbonate makeup annual replacement costs.
Other operating costs include utilities, direct labor, supervision,
maintenance, overheads, and solids disposal. These are shown as a function
of energy utilization and sulfur content of the fuel.
D.
Process Developmental Status
1.
Classification
One difficulty encountered in the cost evaluation of S02 removal processes,
is that the processes are at various stages of development. Table IV - 4 from
Reference 24 classifies most of the known S02 removal processes that are
under development. The classification criteria at the bottom of the table were
established in an American Petroleum Institute study and reported in 1969.
The ordering of the processes within a given classification is random.
The dry limestone, wet limestone, and the Monsanto catalytic
oxidation processes are classified as "first generation." The magnesium
oxide process under development by Chemico has been classified as "advanced
second generation." The modified chamber process developed by Tyco has
been also classified here because advanced pilot plant studies are presently
underway. The molten carbonate process is classified as "second generation"
because only bench studies and conceptual testing have been performed thus
far by Atomics International. Therefore, cost comparisons between these
processes run the risk of being based on nonuniform performance and eco-
nomic data sources.
-------
IV- 44
TABLE IV-4. DEVELOPMENTAL CLASSIFICATION
OF S02 REMOVAL PROCESSES
1.
Company
First Generation
Combustion Engineering
Dry Dolomite Injection
Wet Scrubbing
Catalytic Oxidation
Alkalyzed Alumina
Dry Ab- Adsorption
Dry Dolomite Injection
Dry Removal
Process
Monsanto
U. S. Bureau of Mines
Reinluft
Several
2.
Advanced Second Generation
B&W-Esso
Mitsubishi
Wellman Lord
Kiyoura
Showa Denko
Firma Carl Still
Lurgi
Hitachi
Ovit ron
Chemico
Tyco
Dry Adsorption & Regeneration
Dap Manganese Oxide
Potassium Sulfite
Catalytic Oxidation
Ammonia Scrubbing
Lignite Ash Still
Wet Char Sulfacid
Activated Carbon Adsorption
Wet Scrub-High Acceleration
Magnesium Oxide Scrub
Modified Chamber
3.
Second Generation
A. J. Teller
J. Sieth and Siemens
USBM- Mitsubishi
USBM
Chevron Research
Princeton Chern Research
Ionic s Inc.
Teller Chromatographic
Ferrous Oxide
Red Mud
Phosphate Rock
Catalytic Reduction
Catalytic Requction
Alkaline Scrubbing with
Electrolytic Cells
Manganese Oxyhydride
Molten Carbonate
Potassium Formate Scrub
Mitsubishi
Atomics International
Consolidation Coal
Classification Criteria
1.
First Generation
a. Advanced Pilot Plant Studies
b. Active Research and Development
c. Adaptability to U. S. Market
d. Available Data on Economic Assessment
II.
Advanced Second Generation
a. Meeting First Generation Criteria Except:
1. Less Advanced Pilot Studies or .
Insufficient Information
2. Better Adaptation to Foreign Operations
III.
Second Generation
a. Essentially Bench Studies of New Concepts
b. Pilot Studies Incomplete
c. Economic Assessment Difficult
/1
-------
IV - 45
2.
Problem Areas
Some of the problem areas that are unique for each given process are
identified as a result of the current course of technological development.
a. Dry Limestone. The dry limestone injection technique has ~een
studied intensely. It exhibits the disadvantage of being able to remove under
50 percent of the sulfur oxides contained in the flue gas. Current TV A test
data indicates that at 2.0 stoichiometry, the. S02 removal efficiency varies
between 18 and 26 percent :1:12 percent (Ref. 8).
Besides the problem of low removal efficiency, there are several
other potential problem areas, none of which have been sufficiently studied
to quantify their effects. The addition of limestone into the combustion zone
may adversely affect the combustion process, especially at high stoichiometric
ratios. Increased slagging caused by the effect of CaO on the fusability of
coal ash may inhibit heat transfer. Increased erosion within the boiler may
increase maintenance costs. Electrostatic precipitation may not operate as
efficiently due to the elimination of S03 from the flue gas with subsequent
increase in the average particle resistivity. As resistivity is increased,
the precipitation rate rate decreases. Also as stated earlier, the addition of
limestone base compounds in the ash is likely to increase the problems asso-
ciated with ash disposal. Different methods of disposal may have to be insti-
tuted or existing ones engineered to eliminate the possibility of water pollution.
b. Wet Limestone. Contacting the flue gas with a limestone slurry
significantly increases S02 removal efficiency. The major problem of the dry
limestone process is thus overcome. Efficiency in excess of 85 percent can
be achieved today with prospects of improvement to 95 percent and beyond.
The problems of boiler erosion and reactant disposal identified in the case of
the dry limestone process must also be considered here. However, since the
stoichiometric ratios required approach unity, the magnitude of these problems
are decreased. Since no electrostatic precipitator is required, the problem
of lower precipitator effectiveness is eliminated. The wet limestone process,
however, has several unique problem areas which must be investigated. Cor-
rosion and scaling within the scrubber are major problems which must be over-
come. Engineers are confident that the study of slurry chemistry with respect
to equilibrium com position, pH, degree of supersaturation, rate of desuper-
-------
IV - 46
saturation, and the role played by precipitated salts will result in satisfactory
methods which will control both the corrosion and scaling with the scrubber
and associated components.
c. Magnesium Oxide. Several problem areas were discussed in
Reference 16 concerned with the pilot demonstration performed by Babcock
and Wilcox. Crystalline deposition of magnesium sulfite hydrate in piping,
pumps, sumps, and spray nozzles was a recurring problem. It was believed
to result from the neutralization reaction between bisulfite and magnesia.
Temperature cycling of the scrubbing slurry was also found to contribute to
this deposition. The deposition was prevented by the presence of fly ash
through a scouring action in the absorption slurry. Nonwettable materials
also reduced the tendency to receive deposits. If formed, the deposition can
be removed by simple dissolution.
Deposition of solids where the flue gases first contact the spray
slurry was a problem which occurred in various degrees for both types of
absorbers. These deposits occur near the top of the throat of the venturi
absorber. They also formed at the inlet of the floating bed absorber. These
deposits were identified as being an aggregate of fly ash and MgS03 .3H20.
d. Catalytic Oxidation. Problem areas associated with the Cat-Ox
process were discussed in Reference 17. Principal problems have to do with
fly ash. The fly ash must be removed from the flue gas with very high effi-
ciency in a hot, dry process. Dust not removed from the entering flue gas
tends to plug the catalyst bed in the converter, contaminate the product acid,
and plug the fiber-bed mist eliminator. Monsanto uses an electrostatic pre-
cipitator (or a combination of cyclone separators followed by a precipitator)
for the hot, dry gas cleaning operation. The rate at which the catalyst bed
plugs depends upon the quantity of dust passing through the precipitator. When
the catalyst bed becomes plugged, it is necessary to remove the catalyst for
cleaning. The cleaning operation requires a shutdown of two to three days and
involves the loss of about 2.5 percent of the catalyst by attrition and breakage.
Monsanto estimates that it will be necessary to clean the catalyst at three-
month intervals if the dust passing to the catalyst bed does not exceed the
quantity allowed for in the design. The required dust collector efficiency for
fly ash from firing of pulverized coal is greater than 99.6 percent, which is
higher than is now being provided commercially in fly ash precipitators.
......
tf-r
-------
IV - 47
Of the dust that passes through the converter, a portion is collected
in the acid in the absorption tower and the remainder is collected with very
high efficiency by the fiber-bed mist eliminator, where the buildup of solids
produces an increase in the resistance to gas flow. Monsanto reports that it
has developed a mist eliminator that can be washed periodically, while on stream,
to remove the accumulated solids.
The amount of maintenance required for the Cat-Ox system will be
acutely dependent on the performance of the electrostatic precipitator. A
small deterioration of the precipitator performance may increase the quantity
of dust entering the converter by a large factor, with a corresponding reduction
in the length of the intervals between shutdowns for cleaning. It will therefore
be necessary to attain high standards of maintenance and operation of the pre-
cipitators.
Cat-Ox systems are provided in multiple, parallel independent
trains, so that part of the system can continue to operate even if one train
should fail. The Cat-Ox reheat system can be bypassed in the event of complete
failure. If, however, part or all of the integrated Cat-Ox system should fail,
it is necessary to reduce the load on the boiler or shut it down. Consequently,
the converter in each train is further divided into parallel modular sections
that can be dampered off to permit catalyst cleaning if this should be necessary
between scheduled shutdowns.
The by-product produced by the Cat-Ox system, 78 percent sulfuric
acid contaminated with a small amount (less than O. 1 percent) of fly ash, has
a disadvantage from an economic standpoint. Unless it can be transported by
barge, its movement to markets is severely limited by the cost of shipment.
This is also true, though in lesser degree, of the more concentrated acids,
but the Cat-Ox product acid is additionally limited in applications by its lower
concentration and its impurities.
e. Modified Chamber. Problem areas pertaining to the Tyco process
were discussed in Reference 19. In this system, the scrubber must perform as
a particulate removal device for sulfuric acid mist and residual fly ash. It is
also an absorber for nitrogen oxides and any unhydrated sulfur trioxide. Major
scrubber design and performance parameters need to be confirmed. Additional
information is needed on the effects of fly ash on the scrubber system.
-------
IV -48
Problem areas of the Tyco process include (1) the complexity of
process control and operation, (2) the corrosiveness of the scrubbing acid,
(3) the contamination of fly ash residue in the major product, sulfuric acid,
(4) the need for extremely high efficiency scrubbing to recover nitrogen oxides,
and (5) the high system pressure drop required to prevent sulfuric acid mist
from escaping into the atmosphere. Process control and operation of the
Tyco process will be difficult due to the critical nitrogen dioxide recycle
requirements and the low concentrations of nitrosyl- sulfuric acid in the system.
Close control will be necessary to prevent excess nitrogen dioxide or nitric
oxide losses in the scrubber as changes occur in flue gas compositions.
Startups have to be quick and smooth since it is not practical to store or re-
plenish large quantities of nitrogen dioxide for startups.
With the use of boiler feedwater to remove excess heat from the
circulating acid and maintain a 2500F temperature, some fluctuation of boiler
performance may be encountered during process startup as the switch is made
from preheating the acid with steam to cooling the acid with boiler feed waters.
Proper instrumentation has been recommended to reduce the problem.
The profitability of the Tyco process will be strongly influenced
by the sales value of the by-product acids, 80 percent H2S04 and 60 percent
HN03' In the operating cost estimates of Appendix B acid unit sales revenues
of $4.80 /ton and $15/ton, respectively, were employed. The most desirable
acids are those with high concentration and purity. Nitric acid produced in
the process has these characteris1i:ics, but the sulfuric acid has less adequate
concentration and purity.
f. Molten Carbonate. Major process and engineering problem areas
still requiring solutions prior to full-scale application were identified in
Reference 21. One of these was the elimination of melt carryover with the
absorber off-gas. This loss of melt is undesirable from an economic stand-
point. It is not acceptable from a corrosion standpoint because the scrubbed
gases are returned to the boiler where the materials of construction are not
likely to be compatible with the melt. Furthermore, the melt carryover may
be emitted to the atmosphere from the boiler stack causing additional pollution
problems.
-------
IV - 49
In the absorber, good SO removal efficiencies occur at high
x
velocities, approximately 25 fps. The optimum liquid-to-gas ratio (total mols
of me lt I total mols of gas) is low at O. 0138 and may not be achievable. The
freezing point of the melt is also not too far below the nominal absorber oper-
ating temperature of 850oF.
The sequence of reduction reactions and the required internal design
of the reducer need to be demonstrated. The effect of the elevated operating
temperature of the reducer, 1500oF, on materials of construction is required.
The design of a vessel with three dissimilar materials in contact with each
other presents design problems because of differential expansion. The two-
zone reducer concept is another area which requires additional study to insure
operability. The mode for introducing the coke, e. g., the addition of coke as
a slurry with the melt, is another area where further test work is required.
A primary problem exists to find fly ash and coke filters capable
of producing a solids cake of low liquid content at the process temperature of
850oF. The "dry cake" discharges carry measurable quantities of melt.
These melt quantities must be minimized by means of a satisfactory dry cake
discharge unit in order to reduce the cost for molten carbonate replacement.
Further studies are required to determine the extent to which the
carbonate components can be economically recovered from the solids filter
cakes. A recovery process has been assumed for the lithium carbonate. It
has been assumed that the potassium and sodium salts are not recovered after
the aqueous lithium carbonate recovery process because:
(1)
The fly ash filter cake contains the KCl which is soluble
along with the other potassium and sodium salts.
(2)
The coke filter cake contains the heavy metals which
are removed along with the soluble potassium and
sodium salts.
The discarded sodium and potassium salts are mixtures of carbonates, sulfites,
sulfates, and sulfides in addition to the potassium chloride. The heavy metals
are nickel and vanadium contained in the original coke.
-------
IV - 5 0
E.
Desulfurization of Fuels
An alternate method of 802 abatement is to desulfurize the fuel before
it is burned. Both oil and coal can be desulfurized. This section deals with
the economics involved in oil desulfurization, both domestic and Caribbean,
as well as with the mechanical cleaning of coal. Mechanical cleaning refers
to the removal of pyrites from the coal which tends to reduce the total sulfur
content of the coal. Coal gasification or other of the more complex coal de-
sulfurization methods were not studied since these possibilities will not be
commercially possible until after the 1970s.
1.
Residual Fuel Oil Desulfurization
The economics of residual fuel oil (RFO) desulfurization lack the pre-
cision that has been achieved with respect to the normal oil processing costs.
In part, this is due to the absence of long-term operating data. Large-scale
desulfurization has emerged as a feasible operation only in recent years.
The prediction of incremental costs due to desulfurization processing is com-
plicated by the large number of independent variables. For example, RFOs
of similar sulfur contents which are desulfurized to the same levels can
exhibit significant differences in processing costs. Reactor size will vary
drastically from resid to resid. Thus, capital investment and catalyst costs
can differ significantly. The rate of hydrogen consumption differs depending
on process parameters and extent of desulfurization. Metal content in the
resid determines catalyst life. Generally, specific desulfurization cost esti-
mates must be based on the characteristics of the feed stock, on the level of
desulfurization require, refinery size, and on regional factors such as con-
struction, labor, utility, and hydrogen costs.
With respect to appreciation of desulfurization cost with time, it is
believed that the savings due to process improvements will outweigh the in-
creased costs of labor and materials for some time to come. Catalysts with
ever- increasing resistance to poisoning by sulfur, nitrogen, metals, coke-
formlng materials, and other bad actors in RFO are being found (Ref. 25). It
is expected that the cost of desulfurization will have a downward trend in the
foreseeable future assuming the present trend in technological improvements
continue s.
-------
IV - 51
One of the most extensive efforts in the evaluation of RFO desulfurization
costs to date was undertaken by Bechtel Corporation for the American Petroleum
Institute (API) (Ref. 26). The study analyzed 14 desulfurization schemes with
respect to a typical Caribbean fuel oil normally containing 2.6 weight-percent
sulfur. The usefulness of the results obtained in this study derives from the
fact that the bulk of the RFO imported into the United States comes from the
Caribbean refineries. For example, in 1966, about 3.4 million barrels from
an import total of 3. 8 million barrels were Caribbean RFO.
The various constraints and considerations adopted by Bechtel in evalu-
ating incremental desulfurization costs are summarized below:
(a)
(b)
(c)
(d)
(e)
(f)
Initial sulfur content of 2. 6 weight-percent
Typical large Caribbean refinery size of 300,000 barrels
per stream day of crude oil
A 57.4 volume percent yield of desulfurized No.6 residual
fuel oil
Initial metal content of 500 ppm
Costs derived from incremental investment and operating
costs; any taxes or transportation costs were not included
Costs include sulfur credit based on a market price of $32
per long ton; sulfur credit had only a minor effect on the
overall incremental costs (maximum credit was about 11
percent of final incremental cost)
In order to provide a means of evaluating desulfurization costs for
Caribbean oils of slightly different sulfur contents, some extrapolation was
applied to the Bechtel cost estimates. Thus, there are three curves corre-
sponding to imported Caribbean oils in the nomograph. However, care should
be taken not to venture too far from the 2.6 weight-percent sulfur baseline.
Desulfurization costs are highly sensitive to the type of processing facilities
and the overall properties of the oil feed stock. Whenever a process involves
hydrogen, for example, a major cost item is the hydrogen feed. This cost
can vary drastically with facility location.
The cost differentials shown in the nomograph are based on the assump-
tion that all of the residual oil must be desulfurized. It may be economically
advantageous to desulfurize only part of the resid. One option presented in
-------
IV - 52
Reference 27 is to desulfurize three-quarters of the resid and sell the other
quarter, which is high in sulfur, to users which have stack gas controls.
Under this scheme, the incremental cost of the desulfurized oil would be 50
percent of that shown on the nomograph. Another plan would concentrate the
sulfur into a very high sulfur coke which would be used by the refinery for pro-
cess heat. With the proper incentives, the refinery could install an efficient
802 removal system. The overall economics of this alternative appears
prom ising for those refineries which are "energy poor, " having no immediate
source of inexpensive fuel. If such alternatives are instituted, then the cost
of desulfurization per barrel obtained from the nomograph must be considered
as being conservatively high.
To provide flexibility in treating capital finance, the nomographs allow
variation in the fixed charge rate. Figure B of Nomograph page I may be used
to determine the capital recover factor. Add to this the recurring costs, such
as insurance and taxes, which results in the fixed charge rate.
The Bechtel Corporation assumed that money would be available at six
percent interest with a payout time of five years. Allowing for about two per-
cent of recurring costs, a fixed charge rate of 25 percent results. Thus, a
typical refinery would fix its prices to recover one-quarter of its capital invest-
ment every year.
The fixed and variable costs of desulfurization are calculated separately
in the nomograph. Variable and fixed costs must be added to obtain the total
operating cost. For instance, if we require that the Caribbean oil (average
sulfur content = 2. 6 weight-percent) be reduced to 0.5 pounds of sulfur per
million Btu' s, the following costs are calculated:
Capital investment
6
449/10 Btu/yr
P9/106 Btu
3.49/106 Btu
14.49/106 Btu
Fixed cost (FCR- 250/0)
Variable cost
Total cost
This is equivalent to about 90 cents per barrel of oil having a heating value of
6
6.3 x 10 Btu per barrel.
-------
IV - 53
Incremental desulfurization costs for domestic RFO are based on results
of a study performed by Arthur G. McKee and Company (Ref. 28). In this study,
incremental costs were evaluated for sulfur levels of 1. 0 weight-percent and
0.5 weight-percent, according to the five individual Petroleum Administration
for Defense (P AD) districts.':' The basic approach in the study was to establish
an "average" refinery within each PAD district. Each average refinery was
characterized with respect to crude properties, product rate, product charac-
teristics, processing schemes, and RFO blending requirements.
A major assumption was made with respect to the blending of the heavy
residuum with desulfurized cutter stock. It was assumed that the desulfurized
cutter stock was already available as part of the overall effort in maintaining
a constant sulfur level in the overall domestic RFO pool. Thus, the incremental
costs reflect only the processing necessary for de sulfurizing the heavy residuum.
In the case of PAD Districts 2 and 3, for example, the lower sulfur content
of the crudes obviates additional desulfurization processing. The 1.0 weight-
percent sulfur level in these districts can be achieved simply by blending with
low sulfur cutter stock. Hence, the incremental desulfurization cost is zero.
To obtain the O~ 5 weight-percent sulfur level, it is necessary to partially desul-
furize the heavy residuum in all five PAD districts.
Table IV -5 summarizes the McKee data. Note that the processing scheme
used produced two products: No.2 and No.6 fuel oils. This complicates the
determination of desulfurization costs. If one assumes that there is a strong
market for the additional No.2 fuel oil, a credit can be applied to the costs
since No.2 oil may run as high as $9 per barrel as compared with No.6 oil
at $2 per barrel. However, any real credit is highly dependent on local mar-
ket conditions. Under ideal conditions, where No.2 oil can be sold at a high
price, the cost of desulfurization (allowing a No.2 oil credit) can be reduced
50 percent or more depending on the PAD district. In constructing the nomo-
graphs, no credit was given for the No.2 oil. This conservative assumption
is partially counteracted by the fact that the McKee study assumed the cutter
stock was made available for blending at no cost. What, in effect, has been
done is to assume that No.2 and No.6 oils are essentially the same and that
all the costing is done on a per-barrel-of-new-product basis.
':'The Department of the Interior continues to employ this identification of
different geographical areas in the United States. The districts are:
I- East, II - Midwest, III - Southern, IV - Rocky Mountain, V - West.
-------
TABLE IV-5. DOMESTIC DESULFURIZATION DATA FOR t-I
-<
UNITED STATES PAD DISTRICTS I
CJ1
(Based on Arthur G. McKee and Company Data) *""
DISTRICT
1 2 3 4 5
Minimum Refinery Capacity (b I d) 70,000 40,000 40,000 10,000 42,500
Crude Characteristics
API Gravity 29.7 35.1 34.0 32.6 27. 1
Sulfur (wt %) 1. 33 0.73 0.73 1. 39 1. 29
Resid Data
Sulfur Without Cutter (wt %) 2.6 1.3 1.3 2.7 2.5
Sulfur With Cutter (wt %) 2.0 1. 1 1.1 2.25 1.9
Old No.6 Production (b I d) 11, 340 4,640 5,600 2, 810 10,631
1 % Sulfur Level
New No.6 Production (bid) 9,350 2,250 15, 470
No.2 Production (bl d) 2, 570 730 4, 100
Total New Products (bid) 6 11, 920 2,980 19,570
Incremental Investment (10 $) 7.43 3.19 10.49
Variable Cost ($/day) 6 >:< 4, 372 2,625 5,627
Cap~tal Investment \fl10 ",Btu/yr) 28.8 49.5 24.7
Vanable Cost (9/10 Btu)'" 5.8 14.0 4.6
0.5% Sulfur Level
New No.6 Production (bid) 7,940 3,710 4,480 1,910 13,220
No.2 Production (bid) 4,310 1,210 1,455 1, 152 6,890
Total New Product (bid) 6 12, 250 4,920 5,935 3,062 20, 110
Incremental Investment (10 $) 8.84 5.19 5.72 3.82 11. 38
Variable Cost ($1 day) 6 -', 5, 527 3, 565 3, 858 3,435 6,507
Capital Investment «(;/10 Btu/yr)'" 33.4 48.8 44.6 57.7 26.2
Variable Cost (9/106Btu»:<. 7.2 11. 5 10.3 17.8 5. 1
,,"
"'See text for definitions and assumptions.
-------
IV - 55
The nomographs plot capital investment in terms of cents per million
Btu per year and variable cost in terms of cents per million Btu. These
parameters were obtained from the basic data as follows:
. . 6 / - (100) II
Capital investment (9/10 Btu yr) - (6. 3)(365)TNP x PF
6 - (100) VC
Variable cost (9/10 Btu) - (6.3) TNP
where:
II
TNP =
PF =
VC =
100 =
6.3 =
365 =
=
incremental investment (dollars)
total new product (barrels per day)
plant factor, assumed to be 0.94
variable cost (dollars per day)
cents in a dollar
million Btu I s in a barrel
days in a year
The fixed and variable costs for desulfurization of domestic oils are
nomographed. Variable and fixed costs must be added to obtain the total incre-
mental cost for desulfurization. For instance, in PAD I if we require a limit
of O. 5 pounds of sulfur per million Btu to be maintained, the following costs
are calculated:
6
32.09/10 Btu/yr
8.09/106 Btu
6.79/106 Btu
14.79/106 Btu
Capital investment
Fixed cost (FCR- 25%)
Total cost
Variable cost
This is equivalent to about 93 cents per barrel of oil having a heating value of
6
6.3 x 10 Btu per barrel.
Most desulfurization processes,including the ones used in the Bechtel
and McKee studies, rely on either thermal or catalytic treatment of RFO.
This leads to separation of lighter end products from the residual oil. The
result is a net loss in the oil's heating value. A typical heating value for No.6
oil is 6.3 x 106 Btu per barrel. Desulfurization to low sulfur values (~O. 5 percent
-------
IV - 56
6
sulfur) can lower the heating value to about 5 x 10 Btu per barrel.
effect is included in the nomograph.
This
2.
Coal Desulfurization
a. Introduction. Sulfur is found in raw coal in two principal forms,
organic and pyritic (FeS2) sulfur. Free sulfur exists in coal, but only in
minute quantities. Sulfate sulfur is also present. It varies considerably due
to weathering of coal which oxidizes small amounts of pyrite. This sulfur in
sulfates normally is considered as part of the organic sulfur in the data.
Sulfur can be removed in many ways. The organic sulfur is chemi-
cally bound in a complex manner to the hydrocarbon compounds. Its removal
requires drastic treatments, many of which also result in removal of the pyritic
sulfur. Solvent extraction, liquefication, and gasification have all been studied
as possible methods of removing some, if not all, of the sulfur. Gasification
appears to offer the best chance for development. However, a commercially
competitive process has yet to be discovered. The advent of improved gasifiers
and the continuing increase in the prices of the established fuels could alter
this situation.
Pyrite occurs in coal in bands, in pockets, or as particles, fine or
coarse, mixed with the coal. Pyrite in banks or pockets or as coarse particles
is relatively easy to separate by known washing techniques. Very fine pyrite
becomes increasingly difficult to separate as its size is reduced below that of
the crushed coal size. Before the bulk of the pyrite can be removed from the
coal, the coal must first be crushed fine enough to release the pyrites. Once
released, the pyrite may be removed by electrostatic, magnetic, froth flotation,
concentrating tables, or bacterial methods. The principal method now being
investigated is the use of concentrating tables. Such tables, which use the dif-
ference in specific gravity of the pyrite to separate it, can remove as much
as 70 percent of the pyrite sulfur (Ref. 29).
The present study was lim ited to the economic analysis of conven-
tional coal cleaning methods which utilize concentrating tables for removing
the bulk of the pyritic sulfur.
b. Coal Cleaning Potential. The removal of sulfur from coal by purely
physical means is an inherently limited method of desulfurization. Removing
-------
IV - 57
70 percent of the pyrite usually results in only a moderate decrease in the total
sulfur content. The extent of total reduction depends largely on the proportional
amounts of pyritic and organic sulfur present in the coal. Typically, total
sulfur reductions of between 20 and 40 percent can be achieved. However, it
is difficult to generalize because of the considerable variations in the pyrite-
organic sulfur ratio and in the washability characteristics of specific coals.
It is known that a limited number of U. S. steam coals can be desul-
furized in an economic fashion by pyrite removal. The question may be asked
whether these coals make up a worthwhile portion of the steam coal market.
Figure IV -13 shows the coal reserves of the United States in terms of coal type
and sulfur content. It should be noted that these data are subject to question
since typifying a coal field's average sulfur content based on several sulfur
measurements can be meaningless because of the extreme variability in sulfur
within anyone given seam. In any case, the data shown are considered a best
estimate. Henceforth, the discussion will be limited to bituminous coal. This
coal type makes up the majority of the coal for steam production.
When bituminous coal is pulverized and subjected to a washing
process, certain amounts of the pyritic sulfur and ash are removed. The
degree of removal is dependent, in part, upon the specific gravity of separation.
This parameter is varied by the design and operation of the concentration tables.
As the specific gravity of separation is decreased, more pyrite and ash are
removed. Figure IV -14 shows the relation between specific gravity and removal.
Note that a substantially higher percentage of ash is removed than sulfur. It must
be kept in mind that the removal of ash represents, in itself, an important
improvement. A reduction in ash results in a saving in coal transportation,
in boiler maintenance, in dust removal equipment, and in ash disposal. Thus,
when studying the economics of desulfurization, a credit for ash removal must
be considered which defrays part of the process costs. On the other hand, as
specific gravity of separation is decreased, an increasing percentage of the
combustible material is also removed along with the ash and sulfur. This is
reflected in the "percent recovery" curve in Figure IV -14. In economic terms,
the loss of heating value associated with the combustible material removed is
considered the overriding cost variable which works against the feasibility of
further removal of sulfur or ash. The cost implications of ash removal and
heating value loss will be discussed in the next section.
-------
IV-58
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Figure IV -13. Estimated Remaining Coal Reserves of the United States
. on January 1, 1965
-------
100
(1)
b.O
m
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(1) 40
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IV - 59
80
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. ,..I " , : Ii }\I)" I' '1 .. I, I'" I ! ,I 'i . 1'1 ",.
';:: '-':I~-3'-G- :7~';I~:1:':- I'~' '::-7~l'-:-I'--:;~:'!< ;:i:~'; i (A) Percent Recovery
i/i ~~r :-~ ~~, ;~~'.'-~~ .~:~ !-,.J~!'T~H'/ )~~J':::--ri::'~J
~~:.I, :',~-I.::::,,:i::-=t,.:.~L::'i,~-,~',L~ ,J-:::..-I,.-.: :t::,-,!~'.,!:~~;! (B) Percent Ash Removal.
,..~: -- " :... "I I,,'H, ." , Ii' ',' I,' I" " ',,1,' .1 "I. 'I
-=':' '.1 ~: 'I': [: , ",: ,';, I::' ," , ' n . ::, I . : ' '. "';' " r n Sulfur
..---,' ,... "1-_, -'T" I" h'l,---1'-I, ., "1--..,.- '," '+---'1 (C) Pe ce t Removal
.;; Hi~ ,[Fir '..I.liIi iii, 'lit+ )r,i"J#f'l
"'--, '.. ",':"'r'--, j-"C-"--'-'!.. ,,' .,'","C' "-," ",1,---,--,-,.." ',;,~ '-', " J ,----.,1.".-j
"'I" '1-" If, ',., I,', ",',-- '-lIT 1-" 4111"" "., I,
I rt 'r: ,'11-- ,.J. !t.; -t-;.,, ITJ-" .....,......- ,tT' Tf1~-.1 --I :~ -, ,.~ 'I-J,
"[ !,,' I:: . "',: ," t il :ltJ~t )7-;-,r/ :.':;! :j:ITc, _:': 1 ,":'
:-~::--:-:,-:-:-I-;-,"'-:,,:I:"I'-,'- -c:: lit,1Jt:t'I71,'11I'1', 7',,,' ----"'-1---"',1-,-1:--." NOTES:
,'! ,I "" . ,"i--' -[ !' -'-r-',~I, "I' . -'I" , '.
.::' . ::,: 'T f- --' " " [',' t,-o,' " ,'".., ': '" i 'i
: ii.-:.IF",t:Wl" ~tt-r,O~ T~:,l:41 :,' ~;;;. C, :,l-, ',;:;;', f ~:.-:...:.,..::. r~' '~----F,'::"'T'" ",1---,',---[ These data are based on ave rage
111'j' rl iT',' l!; ut i'!, ::111'-"1" :'" ,1- "1' ',;,::1~ f' ~
~' ~4-:-!,J Ji h':'" '~~"T~I.;: ~-~T~':: .:~~L~-.~~'-~<._,-=-L ..:<.~- values ,of a number 0 bltummous
'.:-!;llh"il'i;:1 .'.''':;! :'I'~: .:.:I'.li,l,:J-;J coal mmes (Refs. 33 and 34).
.,::'i'~-~:,'I:-,.: h):'::j-;-:'-":-IT~ ,,-'r,~::.:!.:.i,': .I-~',,~'II:,J ,:_:i::, ;-:-1
, . , 'I ' L. , , .. .1;,:, . ' " . I' '", t ' ,I '+ ,"',. I.' 1 .', '
,:~~~ :'~~ 1__--~I~1.: ~ ['~:;~I~'~ ;o'~..,i h:'::=.--: ~.--:: 1 --:.:, ::'~-,I,,: ,i - i";o'+
j '1 'I';':, '! (B '1,-,,: ," 'I ..I, 'I'" 'h I--!"" ,,' ",'
:,f ,::' I .', ,':' ." , I;:"" ., ,'[,: u'- ',' I ' ,
"~I:i;' I~-j- -, :' 1":--fJ-:-I-';:I~-':-+':-7':!::;7+.__J 'to'T:--]r
.~ L..1;iJ~';li~ij~It! ,,;:ct~~=i~'..:_I;~~l~~thlj}~[
!..., , i .~' '1' ,u j"'~,"LllJh"'1 ';- I,u ""1--' I". L".I
!-1t --.~..:-(~~{t;Fr': ',I .'h:-I-::t'~~::j~:-: ~':._~'[~"::iT'r~':l-.:t-~'~j
- -----h'" --- L"'-''''i'~--!--'i--:-r- J_-r--T ---I-::-j'.'j
.~~;'~I~:.~F~:ff~I~~t!--lJ-;~~!.Lji~1! .
t,:: t~ h ' ,t:::,; [';, -IT._Lu -I, ,..:1._, -, - [. ,- ,t .j,,' .. J 11 L... ";"
J:J:- _G~__-.L.......j._.- __.L._...~ -_. - ~-- --.. _......~.- ~.:.J..
1.60 1.50 1.40 1.30
Specific Gravity of Separation
60
20
o
Figure IV -14. Effect of Specific Gravity of Separation
on Coal Washing Parameters
-------
IV - 60
The potential of pyrite removal by coal cleaning is dependent upon
the extent of desulfurization possible in the coal fields of interest. References
31 through 35 present data and discussions related to desulfurization potential.
Emphasis is placed on eastern bituminous coal. Figure IV -14 has summarized.
these data in terms of averages. Based on a review of the data, the following
desulfurization schedule was proposed for bituminous coals:
Percent Desulfurization
Applicable % of Coals
o
10
20
30
40
50
>50
o
5
20
50
20
5
o
The schedule assumes that a certain percentage of the United
States bit1,lminous coals can be economically desulfurized by the percentages
given. This schedule was used to estimate the coal cleaning potential of the
known bituminous coals. Figure IV-15 shows the results. Curve A is the
cumulative reserve having a sulfur content of a given percentage of sulfur or
less. Curve B is the coal's sulfur distribution assuming all the coal is cleaned
according to the assumed schedule. The curves may be interpreted as follows:
If a one percent sulfur limit were placed on coal-fired utilities, then there exist
some 216 x 109 short tons of bitumin01,ls coal which may be burned with no addi-
tional desulf1,lrization equipment or coal washing. If the coals having somewhat
higher sulfur contents were processed, then 50 x 109 additional short tons of
coal would become available (the difference between curves B and A). Thus,
on a reserve basis, there would exist some 23 percent more coal than other-
wise would be available at one percent or less. This is the case for the U. S.
as a whole. Of course, if put on an individual coal region or local market
basis, the results would, in some cases, be more promising. Thus, coal
cleaning as a significant and economical general alternative for S02 abatement
is probably limited to specific regions or mines.
c. Coal Desulfurization Economics. A coal processing plant must be
designed for a specific coal. No two coals are identical and thus no one system
-------
800
:j'
.1:
IV - 61
.-- -\:~-~
I
700
600
~-4~
:1 :
:I.:i :
... .1 ..
Ui:;::~~.
,.
: : ~ [ ~:!.l ;: i :
o
1
2
4
5
6
7
Figure IV -15. Sulfur Distribution of Remaining Bituminous
Coal Reserves (1965)
-------
IV - 62
can do an optimum job of desulfurizing them all. The properties of coal
coming from the same seam may vary to such an extent that system efficiency
must be compromised to accept the variability. Sulfur content varies as does
the form in which it is present. Pyrite liberation and washability of the coal
varies. Thus, each installation must be tailor-made (Ref. 37). This factor
introduces scatter in the cost data as do mining costs and variation in coal
sizes. This mustbe kept in mind when typical cost data are presented.
The cost of coal processing was based on two independent studies,
the Paul Weir Company study in 1965 and the Resources of the Future, Inc.,
study in 1968 (Refs. 31 and 32). The general approach of these studies was
to take specific eastern bituminous mines, study their existing mining and pro-
cess methods, and then to upgrade the processing to implement different
degrees of desulfurization. Normally, three levels of desulfurization were
investigated; i. e., those which correspond to separation densities of 1. 60,
1. 45, and 1. 35. Process complexity and cost increase with each step.
Because of the variability in the data, a number of different forms
of data presentation were studied in an effort to minimize cost variability. The
form which appeared to minimize variability made cost dependent upon the per-
cent of desulfurization. Figure IV -16 depicts this relationship based on the
Paul Weir 1965 costs. Considering the vast differences in coal and plant
characteristics, the resulting scatter in cost is reasonably small. Though
not presented, the Resources of the Future, Inc., 1968 cost data for 13 eastern
. mines generally fell within the dashed lines. A comparison of the gro,und rules
utilized in the two studies was made. It was concluded that the Weir study re-
sulted in somewhat higher costs. Thus, the solid line shown in Figure IV-16
was selected as the "typical" cost relation (with the knowledge that it was a
slightly conservative average of the available data). This line was then appre-
ciated based on the Marshall and Stevens Mining Industry Index (Ref. 36). An
appreciation factor of 1. 4 was used to convert 1965 dollars to 1971 dollars.
This appreciation is applicable to both the capital and operating costs. The
fixed and variable costs were then separated and a nomograph constructed.
This nomograph shows capital investment and variable cost as a function of
the run-of-the-mine (ROM) sulfur content and the ultimate sulfur content
required. For example, assume that a coal has a four percent sulfur content
-------
10
8
;:j 6
+->
P::1
!::
0
.~
.......
.......
.~
6
...
a.>
0...
UJ
+->
!::
a.>
U
+->
UJ
0
U
.......
ro
+->
!::
a.>
6
a.>
...
C)
!::
H
IV - 63
Costs Based on a
typical plant size
of 850 tons/hr and,
plant amortized
over 20 years
(Zero interest rate)
Figure IV-l6. Incremental Desulfurization Costs Based On
Five Reference Bituminous Coal Mines
(1965 Dollars)
-------
IV - 64
required and it is de sulfurized 30 percent. Thus, the ROM content is four per-
cent and the ultimate content is 2.8 percent. The capital investment is then
9<; 1106 Btu/yr. Using a fixed charge rate of ten percent per year,
a fixed cost of 0.909/106 Btu results. The variable cost is found to be 4.4<;1
106 Btu. The total operating cost is 5.3<;/106 Btu. This is equivalent to
$1.27/ton for a 12,000 Btu/lb final product.
The above costs do not account for any credits which might be
obtained from the increased heating value and lower ash content of the coal.
In all cases, there is a saving in terms of ash disposal and maintenance costs
associated with the ash. Reference 38 analyzed the cost of removal of ash
from power plant stacks and ash disposal. It was determined that present
rates are about 10<;/ton of coal burned. Reference 39 gives TV A maintenance
costs which are associated with the presence of ash. Average maintenance
costs of about 10<;/ton of coal burned were reported. Since coal processing
typically removes from 25 to 50 percent of the original ash, a credit of 5-10<; I
ton of coal results. Several desulfurization processes keep the as-washed coal
size large enough to allow shipment by train. Based on coal transportation
costs and the known average mileage between mine and utility, a typical trans-
port cost of $3 Iton was assumed. Since the average Btu value of the processed
coal is about five percent greater than unprocessed coal, a saving of about
15<;/ton can be realized. Thus, looking at the extremes, a total credit due to
improvements in the coal lie between 5 and 25<; Iton of coal burned. This is
approximately O. 2 to 1. 0<; I 1 06 Btu. It can be seen that the credits obtained
from the improved coal partially negate the fixed cost of the process. Thus,
one can calculate the desulfurization costs considering only the variable costs
with little loss in accuracy.
F.
Transportation Costs
Some abatement methods require a change in the pattern of fuel trans-
portation or require the shipment of bulk reactants. In an attempt to account
for the added costs resulting from changes or additions in material transpor-
tation, a nomograph has been constructed which allows one to estimate trans-
portation costs.
-------
IV - 65
An example of one possible use of the transportation nomograph is to
analyze the trade-off between using a stack gas S02 control versus western low
sulfur coal. The incremental cost of the stack control process in terms of
mills/Kw-hr can be compared with the premium cost of low sulfur western
coal. The differential costs of the coals at the mine mouth plus the differential
rail costs can be computed resulting in an incremental cost convertible to
mills/Kw-hr. Thus, an economic comparison can be made which would iden-
tify the least expensive control method.
The transportation rates used to construct the cost curves are based on
1968 to 1970 data. They are "average" rates for this time period. As more
efficient methods of transportation begin to dominate a certain transport mode
(e. g., unit trains or super tankers), the average rate should be reduced. Also,
the cost of certain transport modes can vary from day to day or from region
to region because of local factors such as the supply and demand of that partic-
ular mode of transportation. Long-term contracts and special agreements for
shipping large quantities can also result in costs different from the industry's
"average." These market problems are discussed for the more important
modes in the text which follows.
1.
Oil Transportation Costs
The cost of transporting oil involves more variation of costs than any
other fuel. Costs may range from one cent per barrel per 100 miles for long
. haul tankers to 45 cents per barrel per 1 00 miles for short hauls by rail.
a. Oil Transport Costs Via Tanker. The tanker market is a highly
volatile and competitive market. Tonnage is contracted on a voyage basis
("spot" market) or on short-term or long-term charters. The spot market
can fluctuate from day to day over a wide range. If spot prices are kept high
for several weeks, they may have a decided inflationary impact on the charter
prices contracted during the period of high spot prices. Because rates fluc-
tuate so erratically, it was decided to present what is termed the "base rates"
for tanker transport on the nomograph. The current rate structure can then
be factored in using the "World Scale" factor.
Base rates have been placed on all the important runs; e. g.,
Caribbean- United States, Persian Gulf- U. K., or Persian Gulf-Far East runs.
-------
IV - 66
These base rates are equated to a World Scale of 100 (W100). If rates for a
given run happen to be twice the base rate, it would be termed as a rate of
W200.
As an example, the Caribbean- United States run is 2000 miles.
The World Scale base rate at W100 is 25 cents per barrel. In the last half of
1970, spot charter rates varied from W180 to over W320 with an average
near W240 (Ref. 40). Current World Scale rates appear weekly in the Oil and
Gas Journal.
Another important example is the Persian Gulf- U. K. run which is
11,300 miles. This run has significance because it largely influences the rates
of other runs. Its base rate at W100 is $1. 19 per barrel. Actual rates have
varied between W120 and W300 with an average of over W200 for the last part
of 1970. These rates are very high and are deemed unusual for the industry.
In coming years, with the influx of the super tankers into the fleet and the
possible advent of the Suez Canal opening, demand should be lessened and
rates brought back to more "normal" levels. Under normal market conditions,
the super tankers in the 300,000 dwt class would probably operate the Persian
Gulf- U. K. run at below W50 which, no doubt, will substantially affect the total
market price structure (Ref. 41).
b. Oil Transport Via Rail. The cost of rail transportation has
generally decre;:1sed since the beginning of the 1960s. Though it has been
uncommon to sJ;1ip crude or residual oils in tanker cars in the past, this mode
. .
of transportation has been on the increase. Canada has recently begun the first
unit train operation transporting well head oil from its northern provinces.
Such unit trains drastically cut the cost, operating below one-half the normal
25 cents per barrel per 100 mile rate reported for long haul operation greater
than 500 miles. Unlike the ocean tanker market, the rail market is less vola-
tile, being dependent on more long-term influences. Variations in rates are
due mainly to regional cost differences in such items as labor costs, level of
modernization, and local tariff structure. The rates used to construct the
nomograph represent average rates. Rates can vary :1:20 percent. Unit trains
are excluded. Rates for this mode would be approximately one-half the indi-
cated cost per barrel.
-------
IV - 67
c. Oil Transport Via Barge. Barge rates level off beyond 500 miles.
The rates for these long hauls average about 16 cents per barrel per 100 miles.
Significant cost reductions are foreseen in this area as barges carrying 80 to
100 thousand barrels come into use. Long-term charters of such large barges
will significantly affect the cost structure bringing the cost of oil transportation
down to levels as low as five cents per barrel per 100 miles (Ref. 42). Care
should be exercised in using the average barge rate shown on the nomograph
because, as in the case of ocean tankers, rates are highly variable.
d. Oil Transport Via Pipeline. Pipeline transportation costs are
presently averaging about five cents per barrel per 100 miles for long distance
pipelines which operate over 8000 hours per year at capacity (Ref. 42). Costs
skyrocket for short pipelines which have low utilization. As an example, a
7. 6 mile pipeline built by Consumers Power between Edmonton and Bay City,
Michigan, costs the company 46 cents per barrel (Ref. 43). This is equivalent
to a rate of $6. 00 per barrel per 100 miles, which is about 120 times more costly
than the basic long distance rate. Thus, the nomograph which is based on high
utilization pipelines should not be used for short,..run "connect" pipelines which
are subject to considerable variability in cost due to their intermittent operation.
2.
Coal and Bulk Solids Transportation Costs.
a. Bulk Solids Transport Via Truck. Several of the S02 removal pro-
cesses require a supply of rather expensive reactants or catalysts. This ma-
terial will probably be brought by truck. Thus, for convenience in calculating
delivered reactant cost, typical truck shipping rates have been nomographed.
These assume a fully-loa<;led van carrying a minimum shipment of 37,000
pounds (Ref. 10).
b. Coal Transport Via Rail. Coal shipped by rail has normally cost
about one to three cents per ton-mile. The introduction of long haul unit trains
will reduce this rate to about one-half cent per ton-mile. Even lower rates
are projected for utilities requiring large supplies of coal. Such utilities are
planning modern loading-unloading terminals to reduce turnaround time and
will purchase their own unit trains. The nomograph includes two curves, one
for single loads and one for unit train operation (Refs. 10 and 42).
-------
IV - 68
c. Coal ShiI?ments Via Barge. For relatively long hauls (-300 miles),
the TV A estimates barge rates for coal between two and three mills per ton -mile
depending on the annual tonnage of coal contracted. The nomograph presenta-
tion has assumed an annual tonnage of several million tons per year. Adding
a $0. 17 per ton handling cost brings the long haul rate of this annual quantity
to about three mills per ton-mile (Ref. 10).
3.
Gas Transportation
Only one method of transporting gas across the country, pipeline, is
presently used. Long distance rates for pipelines of greater than 500 miles
run between 0.8 and 1. 5 cents per thousand cubic feet per 100 miles (Ref. 42).
The transportation of liquified natural gas by tanker from international
sources is not yet a significant factor in the U. S. fuel supply picture.
-------
V-l
V.
CONCL USIONS AND RECOMMENDATIONS
The major conclusions and recommendations resulting from the con-
tract activities are summarized below:
(1)
(2)
(3 )
(4)
S02 process economics can be nomographed with reason-
able accuracy. In most cases, a nomograph of a process
can be limited to three or four basic cost elements. Sec-
ondary economic variables such as labor rates and utility
rates can be assumed constant with little effect on overall
accuracy.
The parametric study which was performed using the pro-
cess nomographs exemplify the flexibility and usefulness of
the nomographic method. They show that the nomographs
can predict control costs for nearly any set of input param-
eters.
Several nomograph formats were considered during the
course of the contract. The format which was initially se-
lected placed an entire set of curves on a single foldout. This
format resulted in a compact group of nomographs which could
be utilized conveniently with a limited amount of page turning.
An alternate format which was finally chosen placed curves on
a single page. This method though requiring more page mani-
pulations is desirable for nomograph applications since it provides
curves of greater accuracy and allows for duplication of curves
of special interest.
Nearly all cost estimates utilized in the study were based on
dissimilar ground rules. Data presentation formats dif-
fered. Assumptions as to what costs were and were not
chargeable to the abatement process differed. The method-
ology of charging peripheral cost elements such as contrac-
tor fees, indirect field costs, and contingency varied signifi-
cantly. Some estimates were complete while others gave little
backup information. Thus, it is strongly recommended, based
on the difficulties faced in evaluating current economic data,
-------
V-2
that a uniform cost estimate guideline be prepared specif-
ically for S02 stack gas processes. Such a guideline should
include a uniform format for cost presentation. It should
define what is and what is not chargeable to the process.
Also. it should prescribe what data must be submitted as
backup to the cost estimate.
(5)
The nomographs as presented in this report are based on
the best currently available economic data. It is expected.
however. that these data will have to be modified as process
development continues. Prototype tests currently being
conducted or in the planning stages are expected to make
significant contributions to economic data base. Improve-
ments in. or modifications to. system components are also
expected to affect economics. Thus. it is recommended that
the data be reviewed periodically and the nomographs l,lpdated
when necessary. Also. those systems which have not been
nomographed but which are likely to reach commercialization
should be studied and nomographs constructed. This would
include the Chemico version of the magnesium oxide process
which differs substantially from the Babcock and Wilcox process
which was analyzed in the current study.
(6)
Current EP A activities involving the development and demonstra-
tion of S02 control processes should be reviewed with the objective
of assuring that any economic data generated by their activities
be made available. Emphasis should be placed on those cost
elements which are not well known (such as maintenance and
'wear-out of components. waste disposal. or additive consumption).
-------
1.
2.
3.
4.
5.
6.
7.
8.
9.
10.
11.
12.
13.
14.
15.
16.
VI-1
VI.
REFERENCES
Letter dated August 13, 1971 from Louis Scotti, M. W. Kellogg
Co., to Charles Jedlicka, Rittman Associates, Inc.
Letter dated August 18, 1971 from A. G. Sliger, M. W. Kellogg Co.,
to J. P. Dekany, Environmental Protection Agency, under
Contract No. CPA 70-68.
Engineering Science and Technology, June 1970.
Chemical Engineering, April 27, 1970
"Sulfur Oxide Removal from Power Plant Stock Gas, Sorption by
Limestone or Lime, Dry Process, " Tennessee Valley Authority,
1968.
"S02 Removal by Limestone Injection, " Chemical Engineering
Progress Vol. 65, No. 12, December 1969.
"S02 Pickup by Limestone and Dolomite Particles in Flue Gas, "
Journal of Engineering for Power, April 1970.
Telecon with Mr. Richard Stern, EPA Durham, August 19, 1971.
Engineering Science and Technology, June 1970.
"Sulfur Oxide Removal from Power Plant Stack Gas, Use of Lime-
stone in Wet-Scrubbing Process," Tennessee Valley Authority, 1969.
Telecon with Mr. Frank Princiotta, EP A Durham, August 19, 1971.
"Economic Factor in Recovery of Sulfur Dioxide from Power Plant
Stack Gas, " Journal of the Air Pollution Control Association,
January 1971.
"S02 Removal System, Development Proceeds at a Quickening
Pace," Electrical World, May 15, 1971.
Shah, 1. S., "Removing S02 and Acid Mist with Venturi Scrubbers, "
Chemical Engineering Progress, Vol. 67, No.5, 1971, pp. 51-56.
Carlson, G. E. and D. E. James, "Conceptualized Fly-Ash and
Sulfur Dioxide Scrubbing System with By-Product Recovery, " The
Babcock and Wilcox Company Power Generation Division,
January 29, 1971.
Downs, W. and A. J. Kubasco, "Magnesia Base Wet Scrubbing of
Pulverized Coal Generated Flue Gas-Pilot Demonstration, " Re-
search Center Report 5153, The Babcock and Wilcox Company,
September 28, 1970.
-------
VI-2
17.
18.
19.
20.
21.
22.
23.
24.
25.
26.
27.
28.
REFERENCES (Continued)
Semrau, K. T., "Feasibility of New Sulfur Oxide Control Process
for Application to Smelters and Power Plants, Part III: The Monsanto
Cat-Ox Process for Application to Power Plant Flue Gases, "
Stanford Research Institute Final Report, Contract CPA 22-69-78,
1970.
Gruber, A. and A. Walitt, "Oxidation of S02, " Final Report under
Contract No. CPA 70-59, Tyco Laboratories, Inc., October 1970.
"Sulfur Oxide Removal from Power Plant Stack Gas-Tyco Process, "
Draft Report of Preliminary Conceptual Design and Cost Study
under Contract No. TV-29233A, Tennessee Valley Authority,
November 1970.
Tomany, J. P., et aI, "A Survey of Nitrogen-Oxides Control Tech-
nology and the Development of a Low N02 Emissions Combustor, "
A. S. M. E. Paper No. 70-W A/Pwr-2, December 1970.
Drobot, W., S. Finkler and D. R. Whitlock," An Evaluation of
the Atomics International Molten Carbonate Process, " Singmaster
and Breyer, November 1970.
"Development of a Molten Carbonate Process for Removal of Sulfur
Dioxide from Power Plant Stack Gases, Progress Report No.2
(April 1 - October 27, 1968), Part IV, Contractor Development,"
Atomics International Report AI -70- 8.
Cutler, A. J. B., "The Role of Chloride in the Corrosion Caused by
Flue Gases and Their Deposits, " A. S. M. E. Paper No. 70-WA/CD-1,
December 1970.
Wiedersum, G. C., "Control of Power Plant Emissions, " Chemical
Engineering Progress, Vol. 66, No. 11, November 1970, pp. 49-55.
Radford, H. D. and R. G. Rigg, "New Way to Desulfurize Resids, "
Hydrocarbon Processes, November 1970, pp. 187-191.
Cortelyon, C. G., R. C. Mallatt, and H. H. Meridith, Jr., "A
New Look at Desulfurization, " Chemical Engineering Progress,
Vol. 64, No.1, January 1968, pp. 53-59.
"Keeping Sulfur Out of the Stack, " Chemical Engineering/Deskbook
Issue, April 27, 1970.
Sledjeski, E. W. and R. E. Maples, "How Residual Sulfur Limits
Affect Refining," 2 Parts, Oil and Gas Journal, April 29, 1968,
pp. 55-63 and May 13, 1968, pp. 90-95.
-------
29.
30.
31.
32.
33.
34.
35.
36.
37.
38.
39.
40.
41.
42. .
43.
VI-3
REFERENCES (Continued)
Decarlo, Perry, Harry, and Joseph, "The Search for Low Sulfur
Coal, " Mechanical Engineering, April 1967.
"Sulfur Content of United States Coals, " Bureau of Mines IC 8312.
Frankel, Richard J., "Economic Impact of Air and Water Pollution
on Coal Preparation, " Mining Congress Journal, October 1968.
"An Economic Feasibility Study of Coal Desulfurization, " Paul Weir
Company, PH-86-65-29, October 1965.
"A Study on Design and Cost Analysis of a Prototype Coal Cleaning
Plant, " McNally Pittsburgh Manufacturing Corporation, APTD0606,
November 1969.
"An Evaluation of Coal Cleaning Processes and Techniques for Re-
moving Pyritic Sulfur from Fine Coak, " Bituminous Coal Research,
Inc., PH-86-67-139, February 1970.
Tiernan, J. W., "The Sulfur Problem in Coal: What's Being Done
to Get Rid of It, " presented at the Industrial Coal Conference,
University of Kentucky, April 12-13, 1967.
Popper, Herbert, ed., Modern Cost-Engineering Techniques,
McGraw-Hill, New York, New York, 1970.
Brennan, Peter J., "Coal Researchers are Grappling with Sulfur, "
Chemical Engineering, October 9, 1967.
O'Connor, John R., and Joseph F. Citarella, "An Air Pollution Con-
trol Cost Study of the Steam-Electric Power Generating Industry, "
Journal of the APCA, Vol. 20, No.5, May 1970, pp. 283-288.
Yeager, Kurt, and Laurence Hoffman, "The Physical Desulfurization
of Coal - Major Considerations for S02 Emission Control, "
Mitre Corporation.
Van Dyke, L. F., "Tanker Rates will Remain Sky-High through
Winter, " Oil and Gas Journal, November 2, 1970, pp. 43-45.
"Supervessels: A New Navy for Oil" National Petroleum News,
June 1970, pp. 84-86.
"Fuels, " Power, June 1968, pp. 5-46.
Electrical Week, April 12, 1971. p. 4.
-------
A-l
APPENDIX A
NOMOGRAPH EXAMPLES
(Refer to respective nomographs)
-------
A-2
B.
D.
E.
F.
TABLE A-1. PLANT SIZE AND FUEL CONSUMPTION
NOMOGRAPH SAMPLE CASE
A.
Heat Rate
Plant efficiency
Heat rate, 103 Btu/Kw-hr
Capital Recovery Factor
Interest rate, 0/0
Years amortized
Capital recovery factor
C.
Energy Utilization
Net generating capacity, Mw
Plant load factor, % 6
Annual energy production, 10 KW-hr/yr
Plant heat rate, Btu/Kw-hr
Energy utilization, 1012 Btu/yr
Fuel Cost and Consumption
Energy utilization, 1012 Btu/yr
Heating value of oil 7
Annual fuel consumption, 10 bbl/yr
Fuel cost, $ /bbl 6
Energy cost, C; / 10 Btu
Sulfur Content Unit Conversion
Weight-percent of sulfur in fuel
Heating value, oil 6
Sulfur content, lb/10 Btu
Flue Gas Flow Rate
Net generating capacity, Mw
Plant heat rate, Btu/Kw-hr
Flow rate, 106 scfm
(enter)
(read)
(enter)
(enter)
(read)
(enter)
(enter)
(read)
(enter)
(read)
(enter)
(enter)
( re a d)
(enter)
(read)
(enter)
(enter)
(read)
(enter)
(enter)
(re a d)
38
- 9000
8
15
11.5
500
75
3.25
9000
29.6
29.6
0.46
3.5
55.0
3.5
1. 75
500
9000
0.9
-------
TABLE A-2. DRY LIMESTONE
NOMOGRAPH SAMPLE CASE
Net gen~rating capacity, Mw
Plant heat rate, Btu/Kw-hr
Plant load factor, 0/0
As-received heating value of coal,
Weight-percent of sulfur in fuel
Sulfur content of fuel, lb/106 Btu
Energy utilization, 1012 Btu/yr
Flow rate, 106 scfm
Btu/lb
A. Efficiency
Stoichiometric ratio
S02 removal efficiency (optimistic), 0/0
B. Limestone Use Factor
Sulfur content, lb/106 Btu
Stoichiometric ratio
Percent purity
Limestone use factor
C. Fixed Cost
6
Flow rate, 10 scfm
Limestone use factor
Capital cost, $106
Fixed charge rate, % /yr
Fixed cost, $106/yr
D. Other Costs
Energy utilization, 1012 Btu/yr
Limestone use factor, lb/l06 Btu
Other costs, 106 $ /yr
E. Fine Grinding Correction
Limestone particle size, microns
Incremental cost of grinding
F. Additive cost
Energy utilization, 1012 Btu/yr
Limestone use factor, lb/106 Btu
Limestone cost, $ /ton
Incremental grinding cost, $/ton
Disposal cost, $/ton
Total reactant cost, $ /ton
Annual additive cost, $/yr
Total Operating Cost = C + D + E + F
= 0.23 + 0.35 + 0.32 + 0
= 0.90 x $106/yr
200
1, 100
50
12,000
2.0
1. 67
9.6
0.44
(enter)
(read)
( enter)
(enter)
(enter)
(read)
(enter)
(enter)
(read)
(enter)
(read)
(enter)
(enter)
(read)
(enter)
(read)
(enter)
(enter)
(enter)
(read)
A-3
2.5
30
1. 67
2.5
90
14.3
0.44
14.3
1. 15
20
0.23
9.6
14.3
0.35
-----
o
9.6
14.3
4.0
o
0.5
4.5
0.32
-------
A-4
TABLE A-3. . WET LIMESTONE
NOMOGRAPH SAMPLE CASE
Net generating capacity, Mw
Plant heat rate, Btu/Kw-hr
Plant load factor, %
As-received heating value of coal,
Weight-percent of sulfur in fuel
Sulfur content of fuel, lb /106 Btu
Energy utilization, 1012 Btu/yr
Flow rate, 106 sefm
Btu/lb
500
9, 000.
75
12, 000
3.5
2.92
29.6
0.9
A. Blank (an overall S02 removal efficiency of 85% is assumed)
B. Limestone Use Factor
Sulfur content, lb/106 Btu
Stoichiometric ratio
Percent purity 6
Limestone use factor, Ib/10 Btu
C. Fixed Cost
6
Flow rate, 10 sefm
Limestone use factor
Capital cost, $106
Fixed charge rate, %/yr
Fixed cost, $106 /yr
D. Other Costs
Ene rgy utili zation, 1012 Btu / yr
Limestone use factor, Ib/106 Btu
Other costs, $106/yr
E. Reheat Correction
Flow rate, 106 sefm
o
Reheat temRerature, F
Credit, $106/yr
F. Additive Cost
Energy utilization, 1012 Btl1/yr
Limestone use factor, Ib/106 Btu
Limestone cost, $ /ton
Disposal cost, $ /ton
Total reactant cost, $ /ton
Annual additive cost, $/yr
Total Operating 80st = C + D - E + F
= 1. 16 + 1. 22 - 0 + 0.90
= 3.28 x $106/yr
(enter)
(enter)
(enter)
(read)
(enter)
(enter)
(read)
(enter)
(read)
(enter)
(enter)
(read)
(enter)
(enter)
(read)
(enter)
(enter)
(enter)
(read)
2.92
1. 30
90
13.5
0.9
13.5
5.8
20
1. 16
29.6
13.5
1. 22
o
29.6
13.5
4.0
0.50
4.50
0.90
-------
TABLE A -4. MAGNESIUM OXIDE
. NOMOGRAPH SAMPLE CASE
Net generating capacity, Mw
Plant heat rate, Btu/Kw-hr
Plant load factor, %
As-received heating value of coal, Btu/lb
Wight-percent of sulfur in fuel
Sulfur content of fuel, lb / 1 06 Btu
Energy utilization, 1012 Btu/yr
J;"low rate, 106 sefm
500
9,000
75
12, 000
3.5
2.92
29.6
0.9
A. Fixed Cost
6
Flow rate, 10 sefm
Sulfur content, lb/106 Btu
Fixed capital cost, $10~
Fixed capital cost, $10
Fixed charge rate, % /yr
Fixed cost, $106/yr
(enter)
(enter)
(read)
(enter)
(enter)
(read)
B. By-Product Acid Credit
. . . . ..- . 12
Energy utilization, 10. Btu / yr
Sulfur content, lb / 106 Btu
Sulfur removed, 10~ tons/yr:
Sulfur removed, 10 tons/yr
98.5% H2S04 quantity, 103 tons/yr
98.5% H2S04 Ltnit cost, $ /ton
98.5% R2S04 credit, $106/yr
(enter)
(enter)
(read)
(enter)
(re ad)
(enter)
(re ad)
C. Raw Materials Costs
Energy utilization, 10~2 Btu/yr
Magnesia quantity, 10. tons/yr
Magnesia unit cost, $/ton
Magnesia cost, $106/yr
(enter)
(read)
(enter)
(read)
D. Other Operating Costs
Energy utilization, 1012 Btu/yr
Sulfur content, lb /106 Btu 6
Other operating costs, $10 /yr
(enter)
(enter)
(re ad)
. .
Total Operating Cost
::; A-B+C+D
::; 3.27 - 1. 30 + 0.0275 + 2.2,*
6
::; 4.24 x $10 /yr
A-5
0.9
2.92
16.35
16.35
20
3.27
29.6
2.92
43.0
43.0
130.
10
1. 30
29.6
0.655
40
0.0275
29.6
2.92
2.24
-------
A-6
TABLE A-5. . CATALYTIC OXIDATION
NOMOG~APH SAMPLE CASE
Net generating capacity, Mw
Plant heat rate, Btu/Kw-hr
Plant load factor, %
As-received heating value of coal,
Weight-percent of sulfur in fuel
Sulfur content of fuel, lb / 1 06 Btu
Energy utilization, 1012 Btu/yr
Flow rate, 106 scfm
Btu/lb
A. Fixed Cost
6
Flow rate, 10 scfm 6
Sulfur content, lb/10 Btu
Fixed capital cost, $10~
Fixed capital cost, $10
Fixed charge rate, %/yr
Fixed cost, $103/yr
B. By-Product Acid Credit
Energy utilization, 1012 Btu/yr
Sulfur content, lb /106 Btu
Sulfur removed, 1 O~ tons / yr
Sulfur removed, 10 ton.s/yr
78% H2S04 quantity, 103 tons/yr
78% H2S04 un.it.cost, $ Iton
78% H2S04 credit, $106/yr
C. Raw Materials Costs
Energy utilization, 1012 Btu/yr
Ash content in fuel, wt-%
V205 catalyst quantity, ft3/yr
V205 catalyst quantity, ft3/yr
V205 unit cost'6$/ft3
V205 cost, $10 /yr
D. Other Operating Costs
Energy utilization, . 1012 Btu/yr
Other operating costs, $106/yr
.. . Total Operating Cost
= A-B+C+D
= 5.2 - 0.75 + 0.017 + 2.32
6
= 6.79 x $10 /yr
500
9,000
75
12,000
3.5
2. 92
29.6
0.9
(enter)
(enter)
(read)
(enter)
( enter)
(read)
(enter)
(ente r) .
(read)
(enter)
(read)
(enter)
(re ad)
(enter)
(enter)
(read)
(enter)
(enter)
(read)
(enter)
(re ad)
0.9
2.92
25.1
25.1
20
5.02
29.6
2.92
40.0
40.0
150
5
0.75
29.6
10
425.0
425.0
40
0.017
29.6
2.32
-------
TABLE A-6. MODIFIED CHAMBER
NOMOGRAPH SAMPLE CASE
Net generating capacity, Mw
Plant heat rate, Btu/Kw-hr
Plant load factor, %
As -received heating value of coal,
Weight -percent of sulfur in fuel
Sulfur content of fuel, Ib/106 Btu
Energy utilization, 1012 Btu/yr
Flow rate, 106 scfm
B tu / lb
A. Fixed Cost
Flow rate, 106 scfm
Fixed capital cost, $106
Fixed capital cost, $106
Fixed charge rate, % /yr
Fixed cost, $106/yr
B. Raw Materials Costs
Energy utilization, 1012 Btu/yr
Granular charcoal unit cost, $ /ft
Catalyst cost, $106/yr
C. Other Operating Costs
Energy utilization, 1012 Btu/yr
. Other operating costs, $106/yr
D. By-Product Acid Credits
Energy utilization, 1012 Btu/yr
Sulfur content of fuel, Ib/106 Btu
80% H2S04 quantity~ 103 tons/yr
Sulfur removed, 10,j tons/yr
80% H2S04 quantity, 103 tons/yr
80% H2S04 unit cost, $/ton
80% H2S04 credit, $106/year
Energy utilization, 1012 Btu/yr
NOx volume content in flue gas, ppm
,60% HN03 quantity, 103 tons/yr
60% HN03 quantity, 103 tons/yr
60% HN03 unit cost, $/ton
60% HN03 credit, $106 /yr
Total Operating Cost
= A+B+C-D
= 3.6+0.475+2.2
6
= 5.33 x $10 /yr
500
9, 000
75
12, 000
3.5
2.92
29.6
0.9
(enter)
(read)
(enter)
(enter)
(read)
(enter)
(enter)
(read)
(enter)
(read)
(enter)
(enter)
(read)
(read)
(enter)
(enter)
(re ad)
(enter)
(enter)
(read)
(enter)
(enter)
(read)
- (0. 71 + .243)
A-7
0.9
18.0
18.0
20
3.6
29.6
15
0.475
29.6
2.2
29.6
2.82
142.0
39.9
142.0
5
0.71
29.6
400
16.2
16.2
15
.243
-------
A-8
TABLE A-7. MOLTEN CARBONATE
NOMOGRAPH SAMPLE CASE
Net generating capacity, Mw
Plant heat rate, Btu/Kw-hr
Plant load factor, %
As-received heating value of coal,
Weigh-percent of sulfur in fuel
Sulfur content of fuel, Ib/106 Btu
Energy utilizat.ion, 1012 Btu/yr
Flow rate, 106 scfm
B tu / lb
A. Fixed Cost
Flow rate, 106 scfm
Sulfur content, Ib/106 Rt.u
Fixed capital cost, $106
Fixed capital cost, $106
Fixed charge rate, % /yr
Fixed cost, $106/yr
B. Other Operating Costs
Energy utilization, 1012 Btu/yr
Sulfur content, Ib/106 Btu
Other operating costs, $106 /yr
C. By-Product Sulfur Credit
Energy utilization, 1012 Btu/yr
Sulfur content, lb /106 Btu
Sulfur removed ~ sulfur by-product,
103 tons/yr
Sulfur removed ~ sulfur by-product,
103 tons/yr
Sulfur unit cost, $ /ton
Sulfur credit, $106/yr
500
9,000
75
12, 000
3.5
2.92
29.6
0.9
(enter)
(enter)
(read)
(enter)
(enter)
(read)
(enter)
(enter)
(re ad)
0.9
2.92
22.2
22.2
20
4.44
29.6
2.92
2.46
(enter) 29.6
(enter) 2.92
(read) 42.0
(enter) 42.0
(enter) 10
(read) 0.42
D. Raw Materials Costs
Sulfur removed, 103 tons/yr
Coke utilization, 103 tons/yr
Coke unit cost, $/ton
Coke cost, $106/yr
Energy utilization, 1012 Btu/yr
Ash content in fuel, wt-%
M2C03 makeup cost (fly ash fiHer), $106/yr
Sulfur removed, 103 tons/yr
M2C03 makeup cost (coke filter), $106/yr
Energy utilization, 1012 Btu/yr
Chlorine content in fuel, wt-%
M2C03 makeup cost (KCI filter), $106 /yr
Total Operating Cost = A + B - C + D
= 4.44 + 2.46 - 0.42 + 0.240 + 0.58 + 0.065 + 0.18
= 7. 55 x $10 b / yr
(enter)
(re a d)
(enter)
(read)
(enter)
(enter)
(read)
(enter)
(re ad)
(enter)
(enter)
(read)
42
58
10
0.58
29.6
10
0.065
42
0.18
29.6
0.05
.240
-------
B-1
APPENDIX B
COST ES TIMA TES
-------
B-2
TABLE B-1. ANNUAL OPERATING COSTS
DRY LIMESTONE PROCESS
NOTE:
Plant size,Mw
Fixed Capital Investment (FC!)
Stream time, hr/yr
350
$3,400,000
7000
DIRECT COST
1. Operating labor (1 man/shift @ $4. 50/hr)
2. Supervision - 15% of Item 1
3. Maintenance labor and materials - 4% of FCI
4. Plant supplies - 15% of Item 3
5. Utilities
a. Cooling water --- GPM at $0. 025/M gal
b. Process water --- GPM at $0. 20/M gal
c. Electric power 4420 Kw at 6.75 mills/Kw-hr
d. Natural gas --- MSCFH at $0. 40/MSCF
e. Steam --- M lb/hr at $0. 45/M lb
Chemicals and raw materials
a. Limestone 716 TPD @ $2/ton
b. Extra coalo:< 1. 05 TPH @ $8/ton
Solids disposal (ex fly ash) - 29 TPH @ $0. 50/ton
Subtotal Direct Cost (ex credits)
Credits
None
Total Direct Cost
6.
6a.
7.
8.
9.
INDIRECT COST
10. Payroll overhead
11. Plant overhead
12. Total Indirect Cost
20% of (1 + 2)
50% of (1 + 2 + 3 + 4)
FIXED COST
13. Capital charges 18.22% of FCI
(includes depreciation, interim replacements,
insurance, taxes, and cost of capital)
TOTAL OPERATING COST
14. Net Production Cost - -
Item.s (9 + 12 + 13)
UNIT PRODUCTION COST
15. Gross
a. mills/Kw-hr
b. $ /ton S not emitted
16. Net
a. mills/Kw-hr
b. $ /ton S not emitted
Items (7 + 12 + 13)
Items (9 + 12 + 13)
>:
-------
3
Note: M = 10
MM = 106
TABLE B-2. ANNUAL OPERATING COSTS
WET LIMESTONE PROCESS
Plant size, Mw
Fixed capital investment
Stream time, hr/yr
1000
$10,800,000
7000
DIRECT COST
1. . Operating labor (2 men/shift @ $4. 50/hr)
2. Supervision - 15% of Item 1
3. Maintenance labor and materials - 4% of FCI
4. Plant supplies - 150/0 of Item 3
5. Utilities
a. Cooling water --- GPM at $0. '025/M gal
b. Process water - --- GPM at $0. 20/M gal
c. Electric power - 25,310 Kw at 6.75 mills/Kw-hr
d. Steam --- M lb/hr at $0. 45/M lb
6. Chern icals and raw materials
a. Limestone 1237 TPD @ $2/ton
b. Extra coal':< 3 TPH @ $8/ton
6a. Solids disposal (ex fly ash) - 57 TPH @ $0. 50/ton
7. Subtotal Direct Cost (ex credits)
8. Credits
None
9. Total Direct Cost
INDIRECT COST
10. Payroll overhead
11. Plant overhead
12. Total Indirect Cost
20% of (1 + 2)
50% of (1 + 2 + 3 + 4)
FIXED COST
13. Capital charges - 18.220/0 of FCI
(includes depreciation, interim replacements,
insurance, taxes, and cost of capital)
TOTAL OPERATING COST
14. Net production cost
Items (9 + 12 +13)
UNIT PRODUCTION COST
15. Gross
a. mills /Kw-hr
b. $/ton S not emitted
16. Net
a.
b.
Items (7 + 12 + 13)
Items (9 + 12 + 13)
mills/Kw-hr
$ / ton S not em itted
':
-------
B-4
TABLE B-3. ANNUAL OPERATING COSTS
MAGNESIUM: OXIDE PROCESS
Plant size, Mw 1000
Fixed capital investment (FCl)1 - $26,800,000
Stream time, hr/yr . 7000
3
Note: M = 10
MM = 106
7.
8.
DIRECT COST
1. Operating labor (4 men/shift @$4. 50/hr)
2. Supervision - 15% of Item 1
3. Maintenance labor and materials - 4% of FCI
4. Plant supplies - 15% of Item 3
5. Utilities
a. Cooling water 4230 gpm at $0. 025/M gal
b. Process water - 616gpm at $0. 20/M gal -
c. Electric power - 22;656 Kw at 6.75 mills/Kw-hr
d. Natural gas 268"MSCFH at $0. 40/MSCF
e. Steam M lb/hr at $0.45 /M lb
Chemicals and raw materials!
a. Magnesia 5 TPD @ $40/ton
b. Petroleum coke - 645 lb/hr (dry) @ $8/ton
c. Limestone 1500 lb/hr @ $2/ton
Subtotal Direct Cost (ex credits)
Credits .
a. Sulfuric acid (98.5%) - 890 TPD @ $12/ton
Total Direct Cost
$
1fx£.
157,680
23,650
1,072,000
160,800
44,420
51,740
1,070,500
750,400
6.
9.
58,330
18,060
10,500
$3,418,080
(3, 115,000)
$ 303,080
INDIRECT COST
10. Payroll overhead
11. ' Plant overhead
12. Total Ind irect Cost
20%)of (1 + 2)
50%)of (1 + 2 + 3 + 4)
$
36,270
707,070
743,340
FIXED COST
13. Capital charges 18.22% of FCI
(Includes depreciation, interim replacements,
insurance, taxes, and cost of capital)
$4,882,960
TOTAL OPERATING COST
14. Net Production Cost
Items (9 + 12 + 13)
$5,929,380
UNIT PRODUCTION COST
15. Gross
a. mills/Kw-hr
b. $/ton S not emitted
16. Net
Items (7 + 12 + 13)
$9,044,380
1. 29
111.1
5,929,380
0.85
71. 3
Items (9 + 12 + 13)
a.
b.
mills /Kw-hr
$ / ton S not emitted
-------
TABLE B-4. ANNUAL OPERATING COSTS
CAT-OX PROCESS
B-5
Plant size, Mw
Fixed Capital Investment (FCI)
Stream time, hr /yr
1000
$43,400,000
7000
3
NOTE: M = 10
MM = 106
DIRECT COST
1. Operating labor (4 men/shift @ $4. 50/hr)
2. Supervision - 15% of Item 1
3. Maintenance labor and materials - 4. 5% of FCI
4. Plant supplies - 150/0 of Item 3
5. Utilities .
a. Cooling water 548 GPM at $0. 025/M gal
b. Process water - --- GPM at $0. 20/M gal
c. Electric power - 17,800 Kw at 6.75 mills/Kw-hr
d. Natural gas --- MSCFH at $0. 40/MSCF
e. Steam --- M lb/hr at $0.45/M lb
.6. Chemicals and raw materials 3 3
a. Sulfuric acid catalyst - 1330 ft /yr @ $38. 23/ft
7. Subtotal Direct Cost (ex credits) [
8. Credits
a. Sulfuric acid (78%) - 1061 TPD @ $6/ton 100% (1,447,250)
b. Stack gas cooling':< - 162 MM Btu/hr@ $0.40/MM Btu/hr (453,600)
9. Total Direct Cost $ 1,424,070
INDIRECT COST
10. Payroll overhead
11. Plant overhead
12. Total Indirect Cost
20% of (1 + 2)
50% of (1 + 2 + 3 + 4)
FIXED COST
13. Capital charges 18. 22%'of FCI
(includes depreciation, interim r~placements,
insurance, taxes, and cost of capital)
TOTAL OPERATING COST
14. Net Production Cost
Items (r9 + 12 + 13)
UNIT PRODUCTION COST
15. Gross
16. a. mills/Kw-hr
b. $/ton S not emitted
16. Net
Items (7 + 12 + 13)
Items (9 + 12 + 13)
a.
mills/Kw-hr
$ / ton S not emitted
b.
':
-------
B-6
TABLE B-5. ANNVAL OPERATING COSTS
MODIFIED CHAMBER PROCESS
Plant size, Mw ~
Fixed Capital Investment (FCI)
Stream time, hr/yr
1000
$29,800,000
7000
DIRECT COST
1. Operating labor (5 men/ shift @ $4.50 /hr)
2. Supervision - 15% of Item 1
3. Maintenance labor and materials - 5% of FCI
4. Plant supplies - 15% of Item 3
5. Utilities
a. Cooling water 1720 GPM at $0. 025/M gal
b. Process water - 8 GPM at $0. 20/M gal
c. Electric power - 24,730 Kw at 6.75 mills/Kw-hr
d. Natural gas --- \MSCFH at $0. 40/MSCF
e. Steam ---Mlb/hrat$0.45/Mlb
6. Chemicals and raw materials 3 3
a. Activated carbon':' - 68/000 ft /yr @ $13.50/ft
b. Limestone 1. 25 TPH @ $2/ton
I
6a. Solids disposal (ex fly ash) - 1 TPH @ $0. 50/ton
7. Subtotal Direct Cost (ex credit/?)
8. Credits .
a. Sulfuric acid (80%) - 971 TPD @ $6/ton 1000/0
b. Nitric acid (60%) - 148 TPD @ $25/ton 1000/0
c. BFW heating - 170 MM Btu/hr @ $0. 40/MM Btu
9. Total Direct Cost
INDIRECT COST
10. Payroll overhead
11. Plant overhead
12. Total indirect cost
,
200/0=of (1 + 2)
500/0 . 0 f (1 + 2 + 3 + 4)
FIXED COST
13. Capital charges 18.2.2% of FCI
(includes depreciation, interim replacements,
insurance, taxes, and cost of -Capital)
TOT AL OPERATING COST
14. Net Production Cost
(
Items (9 + 12 + 13)
UNIT PRODUCTION COST
15. Gross
a. mills/Kw-hr
b. $/ton S not emitted
16. Net
a.
b.
\
Items (7 + 12 + 13)
)
Items (9 + 12 + 13)
mills/Kw-hr
$ / ton S not emitted
':'Assume one year activated carbon life in catalytic stripper
3
NOTE: M=10
MM=106
$
$/yr
197,100
29,570
1,490,000
223,500
18,060
1,340
1,168,480
918,000
17,500
3,500
$ 4,067,050
(1,359 500)
(648; 0(0)
(474,900)
$ 1,584,650
45,330
970,090
1,015,420
$ 5,429,560
$ 8,029,630
$10,512,030
1. 50
137.7
$ 8,029,630
1. 15
104.7
-------
TABLE B-6. ANNUAL QPERATING COSTS
MOLTEN CARBONATE PROCESS
Note:
Plant size, Mw
Fixed Capital Investment (FCr)
Stream time, hr /yr
1000
$34,600,000
7000
DIRECT COST
1. Operating labor (5 men/shift @ $4. 50/hr
2. Supervision - 15% of Item 1
3. Maintenance labor and materials - 5% of FCI
4. Plant supplies - 15% of Item 3 !
5. Utilities
a. Cooling water
b. Process water
c. Electric power
d. Natural gas
e. Steam
6. Chemicals and raw materials
a. Lithium carbonate - 115 lb/hr @ $0. 42/lb
b. Sodium carbonate - 990 lb/hr @ $0. 02/lb
c. Potassium carbonate - 1684 ib/hr @ $0. 05/lb
d. Petroleum coke - 15.75 TPH @ $8/ton
6a. Solids disposal (ex fly ash) - 1 TPH @ $0. 50/ton
7. Subtotal Direct Cost (ex credits)
8. Credits
a. Sulfur
9. Total Direct Cost
650 GPM at $0. 025/M gal
--- GP~ at $0. 20/M gal
14,575 Kw at 6.75 mills/Kw-hr
--- MSCFH at $0. 40/MSCF
M lb/hr at $0. 45/M lb
- 261 TPD @ $20/short ton
INDIRECT COST
10. Payroll overhead
11. Plant overhead
12. Total indirect cost
20% of (1 + 2)
50% of (1 + 2 + 3 + 4)
FIXED COST
13. Capital charges 18.22% 9f FCI
(includes depreciation, interim replacements,
insurance, taxes, and cost of capital)
TOTAL OPERATING COST
14. Net Production Cost
Items (9 + 12 + 13)
)
UNIT PRODUCTION COST
15. Gross
a. mills/Kw-hr
b. $ /ton S not emitted
16. Net
a. mills/Kw-hr
b. $ /ton S not emitted
Items (7 + 12 + 13)
Items (9 + 12 + 13)
B-7
M = 103
MM = 106
$
$/yr
197,100
29,570
1,730,000
259,500
6,830
688,670
338,100
138,600
589,400
882,000
3,500
$ 4,863,270
(1,522,500)
$ 3,340,770
45,330
1,108,090
$ 1,153,420
$ 6,304,120
$10,798,310
$12,320,910
1. 76
148.4
$10,798,310
1. 54
130.1
-------
TABLE B-.7. PARAMETRIC STUDY
DRY LIMESTONE PROCESS -
td
I
Q:)
Percent of Sulfur Removed from Flue Gas
Fixed Capital Investm6nt ($MM)
350 M\\T 350 M'V'l 350 I-fVJ 175 M~v
29- S 3.5% S* 5% S 3.5% S
.0 --
50 50 50 50
3.2 3.4 3.7 2.3
PI.l\NT SIZE
Salable By-Product
By-Product Production (Tons/Day)
By-produtt Credit ($MM/Yr)
"" 4f - , ..:!.~"'''; .,.....,.. '-" .....;
Total Gross Productlon Cost
$/Ton Sulfur &ot Emitted
" . " /T"TH
e.:l...LS 1:\.\\,
% Increase of Power Cost
($ML"v1/Yr)
... ., ~ , -........ -",-"''''''
1.36 1. 72 2.09 1.05
155.8 112.2 95.5 137.0
0.56 0.70 0.85 0.86
8.3 10.4 12.6 12.7
Total Net Production Cost ($MM/Yr)
$!Ton Sulfur Not Emitted
L,1 ill s /1~:\Tv!-r
?3 Inc.!:'ease
of Power Cost
---
l'Jctes:
---
*BASE CASE
By-Product Values
78-80% Sulfuric Acid
98.5% Sulfuric Acid
60% Nitric Acid
Sulfur
$6/Ton 100%
$12/Ton
$25/Ton 100%
$20/Short Ton
Contained
Plant Load Factor - 80% (7,000 Hr. Stream Year)
Electric Povrer Ivianufacturing Cost - 6. 75 Mills/KWH
Plant Location - Cincinnati, Ohio
-------
TABLE B-8. PARAMETRIC STUDY
WET LIMESTONE PROCESS
Salable By-Product
1,000 M~v 1,000 MW '1,000 fvHv 700 !!:..Jl -100 MW
2°, c 3 E;ge S* 5% S 3.5% S 3.5% S
.'0 oJ . - 0
82.5 90 93 90 90
9.6 10.8 12.0 8.0 4.7
- .
P:'l\NT SIZE
Percent of Sulfur Removed from Flue Gas
Fixed Capital Investment ($~~)
By-Product Production (Tons/Day)
By.- p!"oduct Credit ($lVlM/Yr)
Total Gross Production Cost
$/'I'on Sulfur NotEmi tt-ed--'
Hills/KWH
% Increase of Power Cost
( $MM/Yr)
4.01 5.14 ' 6.00 3.76 2.25
9 7-;-]:---'- '- 65.2 51.4 ' 68.2 ~ -21. 3
0.57 0.73 0.86 0.77 ' 0..80
8.4 10.8 12.7 11. 4 11. 9
'l'ot~l Net Production Cost ($MH/Yr)
S/ToD Sulfur Not Emitted
r~I i 1.1 s I K~'VI-t
% Increase of Power Cost
~
-------
TABLE B-9. PARAMETRIC STUDY
MAGNESIUM OXIDE PROCESS
PLANT SIZE
1,000 ~ny
2% S
Percent of Sulfur Removed from Flue Gas
Fixed Capital Investment ($MM)
91
22.3
Salable
By-Product
By-Prqduct Production (Tons/Day)
487
By-Product Credi-~$MM/Y~)--~.
1~70
Total Gross Production Cost
S/Ton Sulfur Not Emitted
r!; 11 s/~f'7tI
_J...o-J-..- _"'~'J.
% Increase of Power Cost
($MM/Yr)
7.86
172.3
1.12
16.6
Total Net Production Cost ($M~/Yr)
$/Ton Sulfur Not Emitted
~, Q
.j.L . ;:;
98.5% HZSO 4
1,289
4.51
10.38
86
1. 48
21.9
5.86
48.6
0.84
, 2 . ,
- ~. t.!:
700 MW
3.5% S
95
20.5
623
td
.1
I-'
o
400 m7
3.5% S
95
14.1
356
2 ~r8. ------Yo 25'
6.86
-117.9
1. 40
20.7
4.63
80.4
0.96
140 2
4.57-
137.6
1. 63
2(.1
3.33
1()().1
1 1 a
..:- . ~.:.. J
17.6
-------
TABLE B-IO. PARAMETRIC STUDY
CAT-OX PROCESS .
PLANT SIZE
Percent of Sulfur Remov~d from Flue Gas
Fixed Capital Investment ($!v~vl)
Salable By-Product
By-Product Production (Tons/Day)
By-Product Credit ($M~/Yr)
Total Gross Production Cost
$/Ton Sulfur Not Emitted
I;Iilfs7kttIl--'-' ~~~
% Increas~ of Power Cost
( $!vI!v1/Yr )
-~-
--~
Tota.l Net Production Cost ($MM/Yr)
$/Ton Sulfur Not Emitted
Mills/'K~~I}I
% Increa.se
of Power Cost
*BASE CASE
1, 000 Miv
2% S
1 , 0 00 Hiv
3.5% S*
. 1, 0 0 0 M~A]
5% S
82.5
42.3
90
43.4
93
~: 4 ~
. '-:i . .L
78%.H2S04
565
1,061
1,552
0.77
1. 45
2.12
12.12 12.48 12.67
295.5 158.0 109.0
1~-4.~~-' ~1 ." 7-8 ~.l . S 1
25.8 26.4 25.8
10.97 10.58 10.10
265.8 133.9 86.9
1. 57 1. 51 1. 44
23.3 22.4 21.3
Notes:
-By-P::.:-oduct Values
78-88% Sulfuric Acid $6/Ton 100% Contained
98.5% Sulfuric Acid $12/Ton
60% Nitric Acid $25/Ton 100%
Sulfur $20/Si1ort Ton
Plant Load Factor - 80% {7,OOO Hr. stream Year)
Electric Power Manufacturins Cost - 6.75 Mi11s/K~m
Plant Location - Cincinnati, Ohio
700 MN
3.5% S
90
33.2
743
1. 01
4 00 M~J
3.5% S
90
20.6
424
0.58
9.56 5.94
173.5 .188.7
1.9$-- -- . -z--.-l: 2
28.9 31.4
8.23
149.3
1. 68
24.9
5.18
16L~.5
1.85
~.- ~
L./ .~
td
I
......
......
-------
TABLE B-l1. PARAMETRICSTUDY
MODIFIED CHAMBER PROCESS
tJ:j
I
~
I:>.:)
1,000 i~1W 1,000 }fl<'] 1,000 HN 700 ~1W 400 H\Ai'
2% S 3 ~o. S* 5'1. S 3.5% S ":? ~9. S
. =>.0 '0 ..J . :> 0
78 87.5 01 87.5 87.5
""-
29.7 29.8 30.3 22.8 15.5
PLk'IT SIZE
Percent of Sulfur Removed from Flue Gas
Fixed Capital Investment ($~~)
Salable By-Product
80~; H 2 SO ~, 60% HN03
431 971 1,446
148 148 148
"0. 67 'j; :~() ~_. 2.02 "
0.65 0.65 0.65
10.45' 10.51 10.68
268.6 137.7 94.1
1.49 1. 50 1. 53
22.1 22.2 22.7
8.84 8.03 7.33
227.1 104.7 64.6
1. 26 1.15 1. 05
18,;7 17.0 15.6
By-Product Production (Tons/Day)- H2S04
- HN03
By- !':t"-GdU€-l~d-it . . ($f.1M/Yr)
- H2S04
- HNO'3
Total Gross Production Cost
SITon Sulfur Not Emitted
Iv! i 2..1 s II:<~'J I1
% Increase of Power Cost
( $HM/Yr)
Total Net Production Cost ($~~M./Yr)
$jTon Sulfur Not Emitted
Mills/Ki'iTH
% I~crease of Power Cost
680
104
388
59
O. 9 5~" ~o . 54 "
0.,45 0.26
7.99 5.32
149.5 174.1
1. 63 1. 90
24.1 28.1
6.26 4.33
117.0 141.6
1. 28 J..54
19.0 22.8
*BASE CASE
Notes:
~~'-=-Product Values
78-80% Sulfuric Acid $6/Ton 100% Contained
98.5% Sulfuric Acid $12/Ton
60% Nitric Acid $25/Ton 100%
Sulfur $20/Short Ton
Plant Load Factor - 80% (7,ODO Hr. Stream Year)
El ,..,c+-,~i C PO'T""~ ~"Aan'lfa("'-!-'lr' nr< ,""OC"t - 6 -/ r::. M~ 11c/KT.T7J
--'" '-- ~.. ._. \0...: ''''''..L. ;."1 ,~ ..." '-' \,.. ..J.,. ::; '-' ...:.; . ....." .l1... ,,-,... VV.-:.1-
Plant Location - Cincinnati, Ohio
-------
TABLEB-12. PARAMETRIC STUDY
MOLTEN CARBONATE PROCESS
PLANT SIZE
Percent of Sulfu:!:' Removed from Flue Gas
Fi.xed Capital Investment ($MJ.'1)
Salable By-Product
By-Product Production (Tons/Day)
By.-product Cred i t ($r1M/Yr)
'rotal Gross Production Cost
$/T"o n~G.TIU1.'-r-Vt- En1'i.tted-~ I
Mi.lls/K\\E
% Increase of Power Cost
($Y.1M/Yr)
-- ---.------"'---.'-'---
Total Net Production Cost ($MM/Yr)
. $/Tou Sulfur Not Emitted
Mills/KWH .
;~ Irtcrease of
Power Cost
* B.l\SB CASE
No tes:
--
By-Product Values
78-80% Sulfuric Acid
98.5% Sulfuric Acid
60% Nitric Acid
Sulfur
$6/Ton 100%
$12/'fon
$25/Ton 100%
$20/Short Ton
1,000 HVl
2% S
91
31.4
1,000 MW
3.5% S*
95
34.6
1,000M'i'i
5% S
96.5
37~0
SULFUR-
144
380
0.84
10.68
u --z-}"4~-5-' .
1.53
22.7
9.84
216.1
1.41
20.9
Contained
Plant Load Factor - 80% (7,000 Hr. Stream Year)
Electric Power Manufar.::turing Cost - 6.75 Hi1ls/KlvE
. Pler.t Location -Cincinnati, Ohio
2 r 1
0.....
1.52
. 12.32
----14-8-";-4
1. 76
26.1
10.8
130 . 1
1. 54
22.8
2.22
\
13.74
...---.-1-14 .0
1.96
29.0
11.53
95.6
1. 65
24.4
700 HW
3.5% S
95
27.2
400 MI'i'
3.5% S
95
18.2
183
1. Oi
104.4
0.609
9.46 6.25
16 2.-8-~~~1&8 . :
. 1.93 2.23
'28.6 3,3.0
8.40
144.5
1.71
25.3
5.64
170.0
2.02
29.9
td
I
......
'"
------- |