APPLICABILITY OF AQUEOUS SOLUTIONS TO THE
REMOVAL OF SO2 FROM FLUE GASES. VOLUME I
L . E . Gres s ingh , et al
Envirogenics ^Company
El Monte ,. Califo rnia
October 1970
NATIONAL TECHNICAL INFORMATION SERVICE
Distributed ... 'to foster, serve
and promote the nation's
economic development
and technological
advancement.'
-------
ENViROGENICS COMPANY
A DIVISION OF
AEROJET-GENERAL CORPORATION
APPLICABILITY OF AQUEOUS SOLUTIONS TO
THE REMOVAL OF S02 FROM FLUE GASES
FINAL REPORT
VOLUME I
PREPARED UNDER CONTRACT PH 86-68-77
SUBMITTED TO
NATIONAL AIR POLLUTION CONTROL ADMINISTRATION
U.S. DEPARTMENT OF HEALTH, EDUCATION, AND: WELFARE
-------
STAHO/WO TITLE PAGE
FORTECHWCALWEPORTI
I! Report No.
APTD-0619
- Applicability of Aqueous Solutions to the Removal of SO,
From Flue Gas Volume 1
r
I. PWMMMt SfiJJSirtion R«pt. No.
Air Pollution Control Department, Envirogenics
Division of Aerojet-General Corporation
El Monte, California
10. PrafMferTMH/Work IMK No.
ITT.
'Name and)
National Air Pollution Control Administration
Cincinnati, Ohio 45227
PH 86-68-77
13. Ty~pY«rftopoft 4 Period Covered
TOpomorWAfeiicy
18.Abstracts ,xhe program included: literature survey; preliminary economic evaluation for
comparative purposes;.selection of candidate processes; laboratory experimentation to
demonstrate, simplify and improve candidate processes; and process simplification and
improvement of each candidate existing process; demonstration of process feasibility of
candidate new orocess; plant-scale evaluation and cost estimates for the candidate proces-
ses to both new and existing power plant facilities and to a new smelter facility. Of the
four candidate processes the Zinc Oxide process was considered to merit further study, bot
in the form of a fluidized bed and in the form of the original Na+ scrubbing process to
the small-scale pilot stage. The three remaining candidate processes (Cominco Exorption,
Amnonia-Hydrazine Exorption, and Mitsubishi Lime) are not considered to be as economicalljy
attractive as the original Johnstone process. A major problem confronting any aqueous
process in which sulfur dioxide is recovered as such is that of oxidation in the scrubbei
Nearly 700 references are listed, with an author index. \ v.^
17. Key Words and Document Analytic* (a). Descriptors
Air pollution control equipment
Scrubbers
Sulfur dioxide
Bibliographies
Expenses
Feasibility
Oxidation
17b. Identrfters/Open^nded Term
17c. COSATI Field/Group 13/02, 07/01
18. Distribution Statement
Unlimited
19. Security Class(Thls Report)
UNCLASSIFIED
UNRI ASSIFIFO
-------
This report was furnished to the Mr
Pollution Control Office by the
Aerojet-General Corporation in ful-
-------
APPLICABILITY OF AQUEOUS SOLUTIONS TO
THE REMOVAL OF SO2 FROM FLUE GASES
FINAL REPORT
VOLUME I
October 1970
by
L. E. Gressingh, A. F. Graefe, F. E. Miller and H. Barber
Prepared under Contract PH 86-68-77 by the
Air Pollution Control Department, Envirogenics,
A Division of Aerojet-General Corporation,
El Monte, California
Submitted to
National Air Pollution Control Administration
U. S. Department of Health, Education, and Welfare
-------
1 October 1970
i
CONTRACT FULFILLMENT STATEMENT
This is Volume I of a final report submitted to the National
Air Pollution Control Administration in partial fulfillment of Contract
No. PH 86-68-77. This report covers the period 29 December 1967
to 1 September 1970.
Aerojet-General Corporation
^a^g-rvc
L. E. Gressingh
Program Manager
Approved:
E. M. Wilson, Manager ^
Air Pollution Control Department
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This page left intentionally blank.
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TABLE OF CONTENTS
Volume One
PART ONE
GENERAL
Page
INTRODUCTION 1
I. SUMMARY AND CONCLUSIONS
A. Volume I
B. Volume II
PART TWO
ASSESSMENT OF AQUEOUS SOLUTION METHODS - PHASE I
INTRODUCTION __ .. 14
I. SUMMARY _ _______ 15
II. TECHNICAL DISCUSSION _
A. Literature Review _ 1 7
B. Theoretical Chemical Equilibria Considerations _
1. Conclusions _____^__________i___^^____________^_____i_i^_i___ 24
2. Comments on the SO, -Removal Efficiency of Specific
Types of Processes _ 24
3. Sample Calculation _ 26
C. Economic Analysis
1. Basis for Analysis
a. General Considerations _ 28
b. Capital Costs _ 30
c. Operating Costs 35
d. Profitability ...... 30
2. Detailed Economic Analysis
a. Fulham -Simon -Carves Process 45
b. Showa-Denko Ammoniacal Process 53
c. Cominco Process _ 62
d. Cominco Exorption Process 70
-------
TABLE OF CONTENTS (Cont'd)
f. Howden-I. C. I. (Cyclic Lime) and Mitsubishi
D.
g.
h.
i.
j.
k.
1.
in.
n.
o.
P-
q-
r.
s.
t.
u.
V.
Simplified Lime Processes
Battersea Process
Magnesium Hydroxide Process
Magnesium Oxide Process
Manganese Oxide Process
Haenisch-Schroeder Process
Wet Thiogen Process
Ozone -Mn Ion and MnSOA Processes
Sulfidme Process
Basic Aluminum Sulfate Process
Ammonia-Hydrazine Process
Ammonia -Hydrazine Exorption Process
Mitsubishi Ammoniacal Liquor Process
Mitsubishi Manganese Oxyhydroxide Process
Mitsubishi Lime Process
Mitsubishi Red Mud Process
Other Processes
3. Process Selection
a. Introduction
b.
c.
d.
e.
Processes Eliminated
Marginal Processes
Candidate Processes
Sulfur Dioxide as a By-Product
Select Process References
87
96
104
no
121
128
137
146
157
165
172
181
138
190
198
203
207
209
213
214
215
MMBM
220
220
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TABLE OF CONTENTS (Cont'd)
PART THREE
LABORATORY EXPERIMENTATION RELATING TO
CANDIDATE PROCESSES - PHASE II
INTRODUC TION
LABORATORY EFFORT
A. The Zinc Oxide Process
1. Johns tone Method
a. Process Description 229
b. Process Reactions 230
c. Process Simplification and Improvement 230
2. Fluidized Bed Method
a. Process Description 238
b. Process Reactions 239
c. Demonstration of Process Feasibility 240
3. Attempted Synthesis of Zinc Pyrosulfite 259
B. The Cominco Exorption Process
1. Process Description 261
2. Process Reactions 261
3. Process Simplification and Improvement
a. Introduction 262
b. Desorption of Sulfur Dioxide 263
C. The Ammonia-Hydrazine Exorption Process
t
1. Introduction 264
2. Process Description 265
3. Process Reactions Z65
4. Experiments Relating to Process Feasibility
a. Absorption Studies 266
b. Regeneration Studies 271
-------
TABLE OF CONTENTS (Cont'd)
D. The Mitsubishi Lime Process
1. Process Description 274
2. Process Reactions 274
3. Process Simplification and Improvement 274
PART FOUR
PRELIMINARY PLANT-SCALE PROCESS EVALUATION
AND COST ESTIMATE - PHASE III
I. INTRODUCTION ___ 276
II. PROCESS DESIGN
A. General 278
B. Stack Gas Reheat 279
C. SO? Recovery 281
III. ECONOMIC ANALYSIS
A. Introduction 281
B. Caoital Cost Estimate 281
C. Operating Cost Estimate 283
D. Profitability 285
IV. PROCESSES EVALUATED
A. Introduction 285
B. Zinc Oxide Process
1. Process Design 287
2. Capital Costs 287
3. Operating Costs 393
4. Profitability '324
C. Liroe Process
1. Introduction 334
2. Process Design 334
3. Capital Costs
335
4. Operating Costs 335
5. Profitability 035
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TABLE OF CONTENTS (Cont'd)
Page
V. RESULTS OF THE PHASE III EVALUATION
A. Capital Costs 335
B. Operating Costs and Profitability 341
PART FIVE
FUTURE WORK 344
PART SIX
BIBLIOGRAPHY 346
LIST OF TABLES
Number
1 Aqueous Solution Sorption Processes for Removing SO? from
Waste Gases 18
2 Equilibria Data for Various Scrubbing Solutions 25
3 Factored Fixed Capital Estimates 33
4 Raw Material Prices 36
5 Fulham-Simon-Carves Process Using Anhydrous Ammonia:
Chemical Requirements & By-Product Yields 49
6 Fulham-Simon-Carves Process: Capital Cost Estimate Summary 51
7 Fulham-Simon-Carves Process: Operating Cost Estimate
Summary 52
8 Showa-Denko Ammoniacal Process: Capital Cost Estimate
Summary 58
9 Showa-Denko Ammoniacal Process: Chemical Requirements &
By-Product Yields 59
10 Showa-Denko Ammoniacal Process: Operating Cost Estimate
Summary 60
11 Cominco Process: Chemical Requirements & By-Product Yields 64
12 Cominco Process: Capital Cost Estimate Summary 68
13 Cominco Process: Operating Cost Estimate Summary 69
-------
LIST OF TABLES (Cont'd)
Page
Cormnco Exorption Process: Chemical Requirements
fa By-Product Yields _ ______ _ 73
15 Conunco Exorption Process: Capital Cost Estimate 75
Summary _ __ _ _
16 Cominco Exorption Process: Operating Cost Estimate
Summary _ _____ _ __ _____ _ ^
17 Zinc Oxide Process: Chemical Requirements &
By-Product Yields ___ _______ 82
18 Zinc Oxide Process: Capital Cost Estimate Summary _ 83
19 Zinc Oxide Process: Operating Cost Estimate Summary _ 85
20 Howden-1. C.I. Process: Chemical Requirements &
By-Product Yields Using Lime _ ____ _ 89
21 Howden-I. C.I. Process: Chemical Requirements fa
By-Product Yields Using Limestone _ 90
22 Howden-I. C.I. Process: Capital Cost Estimate Summary _ 92
23 Howden-I. C. I. Process: Operating Cost Estimate Summary
Using Lirae _ 93
24 Howden-I. C.I. Process: Operating Cost Estimate Summary
Using Limestone 94
25 Batter sea Process: Chemical Requirements fa
By-Product Yields _ 99
26 Batter sea Process: Capital Cost Estimate Summary _ 100
27 Batt-srsea Process: Operating Cost Estimate Summary
Using Thames River Water _ 102
28 Battersea Process: Operating Cost Estimate Summary
Using Neutral Water _ | 103
29 Magnesium Hydroxide Process: Chemical Requirements
& By-Product Yields _ 106
30 Magnesium Hydroxide Process: Capital Cost Estimate
Summa r y __ _ 109
31 Magaesium Hydroxide Process: Operating Cost Estimate
Summary HI
32 Magnesium Oxide Process: Chemical Requirements
& By-Product Yields _ 116
33 Magnesium Oxide Process: Capital Cost Estimate Summary us
34 Magnesium Oxide Process: Operating Cost Estimate
Summa r y _ _ j 2 o
35 Manganese Oxide Process: Chemical Requirements
& Bv-Product Yields _ 12g
-------
LIST OF TABLES (Cont'd)
Number Pijft
36 Manganese Oxide Process: Capital Cost Estimate Summary__ 127
37 Manganese Oxide Process: Operating Cost Estimate Summary 129
38 Haenisch-Schroeder Process: Chemical Requirements
& By-Product Yields | 133
39 Haenisch-Schroeder Process: Capital Cost Estimate Summary 134
40 Haenisch-Schroeder Process: Operating Cost Estimate
Summary
55 Ammonia -Hydrazine Process: Chemical Requirements
& By-Product Yields
136
41 Wet Thiogen Process: Chemical Requirements &
By- Product Yields _ 141
42 Wet Thiogen Process: Capital Cost Estimate Summary _ 144
43 Wet Thiogen Process: Operating Cost Estimate Summary _ 145
44 Ozone-Mn Ion Process: Chemical Requirements
& By-Product Yields _ _____ 148
45 Ozone-Mn Ion Process: Capital Cost Estimate Summary
(36-sec. Reaction Time) _ _____ 151
46 Ozone-Mn Icn Process: Capital Cost Estimate Summary
(88-sec. Reaction Time) _ _____ 152
47 Ozone-Mn Ion Process: Operating Cost Estimate Summary
(36-sec. Reaction Time) _ _____ 153
4g Ozone-Mn Ion Process: Operating Cost Estimate Summary
(88-sec. Reaction Time) _ _ 154
49 Sulfidine Process: Chemical Requirements & By-Product
Yields _ 159
50 Sulfidine Process: Capital Cost Estimate Summary 1"!
51 Sulfidine Process: Operating Cost Estimate Summary _ 163
52 Basic Aluminum Sulfate Process: Chemical Requirements
& By-Product Yields _ 167
53 Basic Aluminum Sulfate Process: Capital Cost Estimate
Summary _ 169
54 Basic Aluminum Sulfate Process: Operating Cost Estimate
Summary _ 171
56 Ammonia -Hydrazine Process: Capital Cost Estimate
Summary _ 178
57 Ammonia -Hydrazine Process: Operating Cost Estimate
Summary _ 179
5g Ammonia -Hydrazine Exorption Process: Chemical
Requirements & By-Product Yields _ , 183
-------
LIST OF TABLES (Cont'd)
Number Page
*
59 Ammonia -Hydrazine Exorption Process: Capital Cost
Estimate Summary _ . _ .
60 Ammonia-Hydrazine Exorption Process: Operating Cost
Estimate Summary
63 Mitsubishi Manganese Oxyhydroxide Process: Operating
Cost Estimate Summary _ __
61 Mitsubishi Manganese Oxyhydroxide Process: Capital
Cost Estimate Summary _ ________*____- — "
62 Mitsubishi Manganese Oxyhydroxide Process: Chemical
Requirements & By-Product Yields __ __________«.
64 Mitsubishi Lime Process: Chemical Requirements &
By-Product Yields _ _____ _ 199
65 Mitsubishi Lime Process: Operating Cost Estimate
Summary _ _____
66 Mitsubishi Red Mud Process: Chemical Requirements &
By-Product Yields _ . _ 204
67 Mitsubishi Red Mud Process: Operating Cost Estimate
Summary _ _______ _
68 Conversion of Solid ZnO (Kadox-15) to ZnSOj. 2 1/2 HgO at
Room Temperature - Experimental Data ________________________ 243
69 Gaseous Flow Rates and Compositions Used in Reactor
Shown in Figure 38 ... 246
70 Reaction of Fluidized ZnO (Kadox-15) with 0. 27 vol-% SO,
in N2 at 35 C and at 50 C; Experimental Data _ 251
71 Gaseous Flow Rates, Compositions and Space Velocities
Used in Reactor Shown in Figure 40 _ 252
72 Sulfur Dioxide Absorption by Selected Solid Absorbents _ 256
73 Bench Scale SO2 Absorption Data _ 269
74 Capital Cost Estimate Summary _ 282
75 Operating Cost Estimate Summary _ 284
76 Zinc Oxide Process: Capital Cost Estimate Summary,
Case 1 _ _ 295
77 Zinc Oxide Process: Capital Cost Estimate Summary,
Case 2 _ _____ _ 296
78 Zinc Oxide Process: Capital Cost Estimate Summary,
Case 3 _ ^ _ 297
79 Zinc Oxide Process: Capital Cost Estimate Summary,
-------
LIST OF TABLES (Cont'd)
Number
80
81
82
83
84
85
86
87
88
89
90
91
92
93
94
95
96
97
98
99
100
101
102
Zinc Oxide Process: Working Capital, Case 1 ______________
Zinc Oxide Process: Working Capital, Case 2
Zinc Oxide Process: Working Capital, Case 3 ________________
Zinc Oxide Process: Working Capital, Case 4
Zinc Oxide Process: Operating Cost Estimate Summary,
Case 1 __________________________________________
Zinc Oxide Process: Operating Cost Estimate Summary,
Case 2
Zinc Oxide Process: Operating Cost Estimate Summary,
Case 3
Zinc Oxide Process: Operating Cost Estimate Summary,
Case 4
Plume Reheat: Operating Cost Estimate Summary, Cases
1 & 2
Plume Reheat: Operating Cost Estimate Summary, Case 3
Plume Reheat: Operating Cost Estimate Summary, Case 4
Sulfuric Acid Plant: Operating Cost Estimate Summary,
Cases 1 & 2 __________________________________________
Sulfuric Acid Plant: Operating Cost Estimate Summary,
Ca.se 3 ________________________________________
Sulfuric Acid Plant: Operating Cost Estimate Summary,
Case 4 .
Zinc Oxide Process: Annual Raw Material Requirements
and Costs, Cases 1 & 2
Zinc Oxide Process: Annual Raw Material Requirements
and Costs, Case 3
Zinc Oxide Process: Annual Raw Material Requirements
and Costs, Case 4
Zinc Oxide Process: Manning Table and Cost, Cases 1 & 2
Zinc Oxide Process: Manning Table and Cost, Case 3
Zinc Oxide Process: Manning Table and Cost, Case 4
Zinc Oxide Process: Annual Utility Requirements and Costs,
Cases 1 & 2
Zinc Oxide Process: Annual Utility Requirements and Costs,
Case 3
Zinc Oxide Process: Annual Utility Requirements and Costs,
Case 4
299
300
301
302
305
306
307
308
309
310
311
312
313
314
315
316
317
318
319
320
321
322
323
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TBTO*'
Number
103 Zinc Oxide Process: Economic Analysis, Case 1 _ . 326
104 Zinc Oxide Process: Economic Analysis, Case 2 _ m 327
105 Zinc Oxide Process: Economic Analysis, Case 3 _ 328
106 Zinc Oxide process: Economic Analysis, Case 4 _ 329
107 Lime Process: Capital Cost Estimate Summary, Case 3 _ 337
108 Lime Process: Working Capital, Case 3 _ _ _ 338
109 Lime Process: Operating Cost Estimate Summary, Case 3_ 339
110 Capital Investment Summary _ __ _ _ _ 340
111 Profitability: Plants Operating at 90% Plant Factor _ 342
112 Profitability: Plants Operating at 70% Plant Factor _ 343
LIST OF FIGURES
1 Fulham-Simon-Carves Process: Flow Diagram _ _ 46
2 Fulham-Simon-Carves Process Using Anhydrous
Ammonia: Profitability _ 54
3 Showa-Denko Ammoniacal Process: Flow Diagram _ 55
4 Showa-Denko Ammoniacal Process: Profitability _ 61
5 Cominco Process: Flow Diagram ^^^^^^^^^^^^^^^^^^^ 63
6 Cominco Process: Flow Diagram (Smelter Gas) _ 66
7 Cominco Process: Profitability _ 71
8 Cominco Exorption Process: Flow Diagram _ 72
9 Cominco Exorption Process: Profitability _ 78
10 Zinc Oxide Process: Flow Diagram _ 80
11 Zinc Oxide Process: Profitability _ 86
12 Howden-I. C. I. (Cyclic Lime) /Mitsubishi Simplified
Lime: Flow Diagram _ 88
13 Battersea Process: Flow Diagram _ 97
14 Magnesium Hydroxide Process: Flow Diagram 105
15 Magnesium Hydroxide Process: Profitability 112
16 Magnesium Oxide Process: Flow Diagram 114
17 Magnesium Oxide Process: Profitability 122
18 Manganese Oxide Process: Flow Diagram 123
19 Manganese Oxide Process: Profitability 130
-------
Number
20 Haenisch-Schroeder Process: Flow Diagram 131
Zl Haenisch-Schroeder Process: Profitability 138
22 Wet Thiogen Process: Flow Diagram 139
23 Ozone-Mn Ion Process: Flow Diagram 147
24 Ozone-Mn Ion Process: Profitability 156
25 Sulfidine Process: Flow Diagram 158
26 Sulfidine Process: Profitability 164
27 Basic Aluminum Sulfate Process: Flow Diagram 166
28 Basic Aluminum Sulfate Process: Profitability 173
29 Ammonia-Hydrazine Process: Flow Diagram 175
30 Ammonia-Hydrazine Process: Profitability 180
31 Ammonia-Hydrazine Exorption Process: Flow Diagram 182
32 Ammonia-Hydrazine Exorption Process: Profitability 187
33 Mitsubishi Manganese Oxyhydroxide: Flow Diagram 191
34 Mitsubishi Manganese Oxyhydroxide: Profitability 197
35 Mitsubishi Lime Process: Profitability 202
36 Comparative Assessment of Aqueous Based Processes for
Removing SO, From Flue Gases - Capital Investment __^^^i^ 210
37 Comparative Assessment of Aqueous Based Processes for
Removing SO- from Flue Gases - Operating Costs 211
38 Apparatus for Absorption of Sulfur Dioxide by Solid Zinc Oxide 241
39 Conversion of Solid ZnO (Kadox-15) to ZnSO3*2-l/2 H,O at
Room Temperature 244
40 Fluidized Bed Reactor System 248
41 Conversion of Solid ZnO (Kadox-15) to ZnSO3« 2-1/2 H,O in a
Fluidized Bed Reactor 250
42 Laboratory Apparatus for Investigating SO? Removal from
Flue Gas _ 267
43 Absorption of SO, by NH., and N^H. Solutions in Packed
Columns 270
44 Mitsubishi Lime Process: Flow Diagram 275
45 Zinc Oxide Process Flow Diagram - Case 1 - 1400 Megawatt
New Power Plant Facility; Case 2 - 1400 Megawatt Existing
Power Plant Facility 288-289
-------
LIST OF FIGURES (Cont'd)
Number
46
47
48
49
50
51
52
Zinc Oxide Process Flow Diagram - Case 3 - 220 Megawatt
Existing Power Plant Facility
Zinc Oxide'Process Flow Diagram - Case 4 - 222,000 SCFM
Smelter Gas Containing 2. 9% SO2 _
Zinc Oxide Process - Case 1 - 1400 Megawatt New Power
Plant Facility. Break-Even Chart with Conversion of SO-
to H2S04
Zinc Oxide Process - Case 2 - 1400 Megawatt Existing Power
Plant Facility. Break-Even Chart with Conversion of SO, to
Zinc Oxide Process - Case 3 - 220 Megawatt Existing Power
Plant Facility, Break-Even Chart with Conversion of SO, to
Zinc Oxide Process - Case 4 - New Smelter Facility. Break-
Even Chart with Conversion of SO2 to H-SO. _____________„___
Lime Process Flow Diagram - Case 3 - 220 Megawatt Existing
Power Plant Facility _
290-29l|
292-293J
330
331
332
333
APPENDIX A.
APPENDDC B.
APPENDDC C-l.
APPENDDC C-2.
APPENDDC C -3.
APPENDDC C-4.
APPENDDC C-5.
APPENDICES
Theoretical Chemical Equilibria Considerations _
Conversion of Gaseous Sulfur Dioxide to Marketable
Products: Cost Estimates _
Acknowledgements _
Zinc Oxide Process, Case 1 and 2, Equipment List
Zinc Oxide Process, Case 3, Equipment List
Zinc Oxide Process, Case 4, Equipment List
Lime Process, Case 3, Equipment List
A-l
B-l
C»l
C-2
C-12
C-22
C-35
-------
PART ONE
GENERAL
I. INTRODUCTION
The initial objective of Contract No. PH 86-68-77 was to assay the fea-
sibility of using aqueous systems for removing sulfur dioxide from flue gases.
The period of service of the initial program was from, 29 December 1967 to
31 May 1969. This technical effort is reported in Volume I, "Applicability of
Aqueous Solutions to the Removal of SO- from Flue Gases. " An extension of
this program covered the period 31 May 1969 to 1 September 1970. This part
of the program is reported in Volume II, "The Development of New and/or
Improved Aqueous Processes for Removing SO, from Flue Gases. "
C»
The general discussion, Part One, consisting of the Introduction, and
Summary and Conclusions, is identical in both volumes of this report and
provides a resume of the entire project.
The following three phases define the program effort of the initial
period:
Phase I. Assessment of Aqueous Solution Methods
• Literature survey,
• Preliminary economic evaluation for comparative purposes ; ,
• Selection of Candidate processes^
Phase II. ~ Laboratory Experimentation Relating toiGandidate Processes'^!" ""
• Process simplification and improvement of each candidate
existing process: ,
• 'Demonstration of process feasibility of any candidate new
-------
f-'' I ''
Phase III. Preliminary Plant-£cale »»«€ess Evaluation and
Cost Estimates for the Candidate Processes —.
• Application of processes, selected onjjhe
* of Phases I and II'to both new and existing power
plant facilities *> ' -,
• Application of processes selected on the basis
of Phase's Tand II io a new smelter facility, ._ ^
Phase I was accomplished during the first five months of 1968 with
Phases II and III conducted concurrently during the remainder Of the
calendar year.
i
The following parts of Volume I are concerned with the result* of
the application of the various tasks listed above. Parts Two to Four
cover the work conducted under Phases I, II, and III, respectively. Part
Five discusses recommendations for future work under the contract ex-
tension. Part Six, "Bibliography" is the result of the extensive literature
• - j ,...»—L. _..—..- -.n^,-.-.•r»«|tf m-'^-ni 1.1 »•!••'• minnmt •, , ,M| ^i
survey which was carried out at the beginning of Phase 1?"^Nearly 700
references are listed, togothe-y with an appropriate author index. ,„,«£•£
The program extension, designated as Phase IV, consisted of the
following tasks:
A. Conceive New Aqueous Scrubbing Processes.
B. Develop Improvements to Previously Conceived Aqueous
Scrubbing Processes.
C. Determine the Degree to Which Inadvertent Sorbent
Oxidation Can Be Minimized.
D. Determine the Degree of Interference Which Inadvertent
Sorption of NO Has On SO, Removal Efficiencies.
X £i
E. Support the Laboratory Investigations with Preliminary
Process Evaluations and Economic Analyses.
The tasks of Phase IV are covered in Parts Two through Six of
-------
II. SUMMARY AND CONCLUSIONS
A. VOLUME I
Approximately 500 technical documents, selected from the
bibliography of Part Six, were collected, catalogued, and reviewed for
the identification and description of various aqueous processes which
have been used, or are currently being investigated, developed, or used
for the removal of sulfur dioxide from flue gases. Some thirty processes
were identified, and of these sufficient data were available for a prelimi-
nary economic evaluation of twenty-two. As a result of the evaluation
the following four processes were considered to merit further investiga-
tion:
• Zinc Oxide Process (Sodium Sulfite Scrubbing)
• Cominco Exorption Process (Ammonium Sulfite
Scrubbing)
• Ammonia-Hydrazine Exorption Process
(Hydrazine Scrubbing)
• Mitsubishi Lime Process (Lime water Scrubbing)
The Ammonia-Hydrazine Exorption process, conceived at Aerojet, repre-
sented a paper study, subject to an experimental demonstration of process
feasibility.
Following process selection, a laboratory program was con-
ducted relating to process demonstration and/or improvement. Attempts
to improve the Zinc Oxide process were mainly concerned with lowering
the calcination temperature required for the release of sulfur dioxide from
the regeneration feed, zinc sulfite. No significant improvements were
effected, but the investigation led to the conception of a new process based
on the use of zinc oxide, in which a fluidized bed of this material is used
directly for the low-temperature (50 C) sorption of sulfur dioxide. The
-------
The Cominco Exorption process suffers from the relatively
high steam costs associated with the desorption of sulfur dioxide from
aqueous ammonium bisulfite solution. The use of acids as promoters for
this reaction wajs therefore investigated. Although several acids were
found to be partially effective it was found that the cost of the acid, the
additional complexity to the process caused by the use of the acid, and
other factors would not be compensated by the limited reduction in steam
requirements which might be attainable in this manner. It was concluded,
therefore, that this process must be considered as uneconomical.
The Ammonia-Hydrazm« Exorption process was designed to
combat the high steam requirements of the Cominco Exorption process
through the use of hydrazine as the absorbent for sulfur dioxide. Since
hydrazine salts are highly soluble in aqueous media it appeared that the
desorption of sulfur dioxide from aqueous hydrazine bisulfite might be
effected without the simultaneous volatilization of large quantities of water.
As the result of an experimental program designed to demonstrate process
feasibility, it was found that concentrated hydrazine sulfite solution readily
absorbed sulfur dioxide under simulated process conditions. However, an
unavoidable loss of hydrazine by oxidation occurred during the regeneration
reaction, so that any savings in steam costs through the use of this method
was nullified. Therefore, the process was no longer regarded as
economically feasible.
No experimental work was indicated relating to the Mitsubishi
Lime process. The process was regarded as economical, provided that
by-product gypsum could be sold in quantity. However, a subsequent mar-
ket survey indicated that gypsum requirements could readily be filled from
natural deposits and that no appreciable synthetic gypsum market exists at
the present time. A simplified version of the process, in which the gypsum
is discarded as waste, appeared more attractive. A laboratory effort
-------
In addition to the laboratory effort described above, which
was designed to overcome problems associated with specific processes,
some attention was also directed to a problem which is common to all
aqueous scrubbing methods in which sulfur dioxide is recovered as such;
namely, the oxidation of sulfite to sulfate in the absorber. The literature
indicates that in some processes oxidation can amount to 10 to 14% or
more (expressed as a percentage of the incoming SO9). It was planned
c*
to investigate the extent to which oxygen and nitrogen oxides in flue gas
contribute to oxidation, and to investigate the use of various oxidation
inhibitors, such as hydroquinine, for its prevention. Work in this area
was initiated toward the end of the contract period, and was completed
during the second year of the program (Phase IV). The results are re-
ported in Volume II.
Of the four candidate processes, only the Zinc Oxide pro-
cess (Na scrubbing/ZnO regenerant) was considered for a complete
evaluation in Phase III. Thus, process evaluations and cost estimates
were completed for large new and existing power plants, a small existing
power plant, and for an existing smelter facility. An evaluation was be-
gun on the simplified lime process, but was not completed since the
analysis of limestone systems was being done on another contract.
The major conclusions which were drawn from the work
reported in Parts Two to Four of Volume I, are the following:
• 'Of the four candidate processes which were selected
for further study as the result of the Phase I effort,
the Zinc Oxide process was considered to merit
further study, both in the form of a fluidized bed
system, as proposed by Aerojet, and in the form
of the original Na .scrubbing process^ as developed
by Johnstone to the small-scale pilot stage. For
the Johnstone process available data in Phase III
indicated that for a large power plant (2. 5 MMSCFM
of flue gas) to be operated at break-even conditions,
product sulfuric ac;d would have to be salable at al-
-------
feed gas concentration of 3000 ppm). If, however,
the product acid from this plant could be sold for
only $10/ton, the operation of this "add-on" SO2
Control process would represent a net cost to the
utility of about $1. 23/ton of coal burned. Applied
to a medium-sized smelter effluent (220, 000 SCFM
of the flue gas), the Johnstone process could be
operated at break-even conditions, if product
"""- - sulfuricacid were sold at about $18/ton.
• /'"The three remaining candidate processes (Cominco
Exorption, Ammonia-Hydrazine Exorption, and
Mitsubishi Lime) are not considered to be as
economically attractive as the, j"ohnstone process. ^
• ~A major problem confronting any aqueous process
in which sulfur dioxide is recovered as such is that
of oxidation in the scrubber, .Such oxidation in-
evitably leads to the formation of sulfate, which ~"
is in general less readily isolated from aqueous -/
solution and less readily decomposed than the (
corresponding sulfite. As a result it may be
anticipated that equipment and operating costs
will increase, and product yields will decrease,
in proportion to the extent of oxidation encountered.
B. VOLUME II
In the area relating to new aqueous processes for the re-
moval of SO2 from flue gas, attention has been focused on the use of
fltiidized solids as absorbents. The absorption step, which is conducted
at 50 to 60 C, requires the presence of appreciable water in the gas
phase, and this is provided largely through the use of an aqueous pre-
scrubber. The prescrubber also serves the function of removing SO
-------
Only basic materials have proved suitable as SO2 absor-
bents. It was found, for example, that both alkali (Na, K) and alkaline
earth (Mg, Ca) sulfites are too weakly basic to absorb, but that
carbonates (Na) and oxides (Zn, Mg) are good absorbers. For example,
when zinc oxide was used it was found that more than 50 g of SO, was
absorbed per 100 g of the oxide before SO, removal efficiency dropped
below 90%.
A problem that arises in all regenerative aqueous SO,
scrubbhg processes is that a portion of the absorbent becomes oxidized
by the O~, fly ash, and/or NO components of flue gas. This is highly
«• Jt
undesirable, inasmuch as the absorbent cannot readily be regenerated
for reuse from the oxidized product. The extent of oxidation appears
to be substantially less for essentially dry fluidized bed absorbents
than for bulk water systems, and in particular fly ash, which tends to
catalyze the oxidation in bulk water systems, was determined to be
without effect when incorporated into a dry fluidized zinc oxide absorber.
Of the three gaseous oxidizing components of flue gas (O,»
NO, and NO-}, it was found that NO- is by far the most active, and that
fluidized zinc oxide was partially converted to sulfate when both SO,
and NO- were present in the influent gas. It was discovered, however,
£t
that oxidation of the bed material could be essentially eliminated through
the incorporation of ferrous ion into the aqueous prescrubber. The main
function of the ferrous ion is considered to be that of reducing NO, to
NO. For an influent gas containing all flue gas components, less than
one-half percent of the SO, absorbed by zinc oxide was converted to
£•
sulfate when the prescrubber contained 1% ferrous sulfate. In the prac-
tical case, ferrous ion would be provided to the prescrubber in the
form of scrap iron.
It may be presumed that the use of fluidized solids as SO,
absorbents will be attended to some extent by the attrition of solid
particles, and consequently son-ie attention was devoted to a study of
both particle size and particle activity when zinc oxide was used as
-------
oxide particles are readily attrited to fine particles, but that if the oxide
is first converted to the sulfite through SO2 absorption, and the sulfite is
then thermally decomposed, the resulting regenerated oxide shows con-
siderable resistance to attrition. Recent studies have indicated that both
zinc sulfite and regenerated zinc oxide in the range of -12 to +24 mesh
can be utilized with considerable resistance to attrition at a superficial
gas velocity of about 3 feet per second. It was also found that the re-
generated oxide is much more active toward SO2 absorption than is the
fresh, unused commercially available oxide.
In the area relating to process improvement, attention was
directed toward the thermal decomposition of metallic sulfites, inasmuch
as these compounds are the principal products resulting from the absorp-
tion of SO, by metallic oxides and hydroxides, and because the regenera-
tion of the oxides from the sulfites is probably best accomplished by
thermal means. The decomposition reaction is always attended to some
extent by the formation of sulfate, and other products, as a result of the
disproportionation of the sulfite.
The thermal decomposition of zinc sulfite was studied in
both muffle and tube furnaces as a function of time and temperature. An
important result of this work was the discovery that the rate of the de-
composition reaction was markedly increased in the presence of steam.
It was subsequently found that the use of steam permitted the decomposi-
tion to be carried out at temperatures («300 C) below which dispropor-
tionation occurs, so that the formation of sulfate and sulfide could be
essentially avoided. In another series of experiments it was found that
the use of steam was much less effective in promoting the decomposition
of zinc sulfate.
The disproportionation of magnesium sulfite is more exten-
sive than that of zinc sulfite, and it might be expected that the use of
steam would be less effective in the case of magnesium. Preliminary
experiments have shown that the temperatures required for the decom-
position of magnesium sulfite are substantially decreased through the
use of steam, but that the formation of appreciable sulfate attends the
decomposition.
-------
As a result of the experimental work considered above relat-
ing to the use of fluidized zinc oxide as an absorbent for SO,! a tentative
system was formulated involving the recovery of the SO, as such. The
overall system involves the use of an aqueous prescrubber, the removal
of SO2 from the water-saturated gas by the oxide, and the thermal decom-
position of the resulting sulfite at 275°C for the regeneration of the oxide
and the recovery of SO,. Any zinc sulfate formed is separately decom-
posed at higher temperatures, and no waste product results. In an
alternative system, sulfate is removed by filtration rather than by cal-
cination. To accomplish this, a portion of the sulfite-sulfate mixture
is dissolved in aqueous SO, and the sulfite is re precipitated with zinc
oxide. After filtering the zinc sulfate solution, the sulfite cake is re-
turned to the process.
The studies on oxidation and oxides of nitrogen in aqueous
solution scrubbing systems were combined due to the contributions of
nitrogen oxides to sorbent oxidation. Most of the experiments were
made with sodium sulfite-bisulfite solutions similar to that used in the
Johns tone Zinc Oxide process. A "once through" counter cur rent absorp-
tion column was used in most of these tests. Fresh absorbent solution
was fed to the top of the column and the spent solution removed from
the bottom. Another arrangement was used for some tests in which the
absorbent was recirculated through the column. Other absorbent sys-
tems checked were potassium sulfite-bisulfite, and magnesium, calcium,
and sodium hydroxide solutions.
Commercially available inhibiting and complex ing agents,
widely used in other applications, were screened for their ability to re-
duce oxidation of sulfur dioxide (sorbent) in the scrubber. Oxidation of
the sorbent due to oxygen or fly ash in the flue gas was suppressed by
some of the materials. When nitrogen oxides were present in the flue
gas, however, oxidation was lowered only by using nitrilotriacetic acid,
and this inhibitor was effective only in a potassium sulfite-bisulfite
-------
It was found that, although oxygen in the flue gas contributes
to the oxidation of the sorbent during scrubbing, the high levels of oxida-
tion was progressively greater as the concentration of nitrogen dioxide
in the flue gas was increased. The rate of oxidation was highest in tests
made with 400 ppm each of nitrogen oxide and nitrogen dioxide.
Fly ash did not significantly increase oxidation in systems
where fly ash free absorbents were fed to the once through column. In
absorbent recirculating systems, in which the fly ash accumulated and
some of the iron content was solubilized, a low level of SO2 oxidation
was experienced. The oxidation increased with increasing turbulence in
the system. A similar effect was found when ferric ions such as Fe2(SO4)3
were added to the system.
Saturating the sodium sulfite-bisulfite scrubbing solution
with sodium sulfate inhibits oxidation. This is explained by the limited
solubility of oxygen in high ionic strength aqueous solutions.
Since oxygen is only slightly soluble in water, the liquid
phase is the limiting resistance to the absorption of the oxygen. Thus,
increasing turbulence in the scrubber improves the absorption of oxygen
and the amount of oxidation of the sorbent increases with the turbulence
of the system.
As discussed in Part Three, a pre scrubber circulating a
solution containing ferrous ion removes the nitrogen dioxide from the
flue gas stream. Using this pre scrubber system in conjunction with
aqueous solution scrubbers also reduced oxidation of the sorbent to a
very low level due to removal of the nitrogen dioxide.
Although additional investigations would be needed to verify
the data, it seems that the absorption of nitrogen oxides simultaneously
with sulfur dioxide is about the same quantity as the percent nitrogen
dioxide in the flue gas. The experiments also indicate that the absorption
of N0x into S02 scrubbing solutions has no effect on SO2 removal efficiency.
-------
Miscellaneous process and economic evaluations were made
on the Johnstone Zinc Oxide process, the new Fluidized Zinc Oxide pro-
cess, and a Magnesium Base Slurry SO2 Scrubbing system.
Evaluations involving the Johnstone Zinc Oxide process
included an analysis in which sulfur dioxide recovered from the absor-
bent was converted to sulfur using the Asarco process. If product sulfur
could be sold for $20 per long ton, the net cost of operating this SO, re-
moval/sulfur recovery process on a 1400 MW power plant (at a 70% load
factor) would approximate $1. 36 per ton of coal burned. The economics
of converting the sulfur dioxide to sulfuric acid (see Volume I) was re-
evaluated on the basis of lower sales prices for the sulfuric acid produced.
An analysis of using reverse osmosis to separate the oxidation product
from the absorbent indicated an uneconomical system based on current
technology.
The evaluation of the optimized new Fluidized Zinc Oxide
process showed relatively low capital and operating costs for a system
serving a 1400 MW power plant; however, it must be recognized that
this projection is based on the presumed validity of data that has been
generated on a very small-scale laboratory equipment.
The cost study of the Magnesia Base Slurry SOg Scrubbing
system was made only on the absorption system. An evaluation of the
regeneration system, which was not available, would have to be made to
complete the analysis.
The major conclusions which have been drawn from the work
reported in Parts Two to Six of Volume II are the following:
• Efficient absorption of SO? at flue gas con-
centrations can be effected through the use
of dry, fluidized basic materials in the range
of 50 to 60 C, if sufficient water is incor-
porated into the gas phase upstream of
sorbent contactor.
-------
The formation of sulfate can be essentially
eliminated in a fluidized bed absorber, and
reduced to a very low value in an aqueous
absorber, through the use of ferrous ion in
an aqueous prescrubber to reduce NO2 to
NO.
The thermal decomposition of both zinc and
magnesium sulfites is markedly promoted by
the presence of steam. The use of steam
permits the decomposition of zinc sulfite to
be carried out at a temperature below that
at which disproportionation occurs.
A new process; for the removal of SC^ from
flue gas is described in which dry fluidized
zinc oxide is used as the absorbent. The
oxide is recovered for reuse upon thermal
decomposition of the resulting sulfite, and
the liberated SO, is recovered as such.
Little or no sulfate is formed.
NO (especially NO~) is the major contributor
3C £*
to oxidation of the sorbent in aqueous solution
systems.
In general, the inhibitors and complexing agents
investigated did not lower the level of oxidation
in the presence of NO in the flue gas.
JL
The level of oxidation is less in sorbent solutions
saturated with an inert salt.
The efficiency of SO^ removal from flue gas is
not affected by the presence of NO .
Jw
The economics of the conceptualized fluidized-
bed zinc oxide process appear to be superior to
other regenerable processes for the removal of
SO2 from flue gases, but the state of development
of this process is in its very early stage.
-------
One, merely tentative, "conclusion" bears mentioning:
• It appears that adding NO- to flue gas to
£»
obtain an equimolar ratio of NO/ NO,
prior to scrubbing the gas with aqueous
sulfite-bisulfite solutions or slurries, for
SO,/NO removal will not lower the NO
£ X X
content of the gas significantly, but will
cause unwanted oxidation of the sulfite to
sulfate to increase drastically.
-------
PART TWO
ASSESSMENT OF AQUEOUS SOLUTION METHODS - PHASE I
*
I. INTRODUCTION
A comprehensive literature survey constituted the first task required in
Phase I. The information generated from the survey provided the basis for
assessment of the many aqueous-based scrubbing processes which have been
developed over a period dating back to the latter part of the 19th Century.
Typical information related to process development application, process
reactions, equipment, by-product recovery, economics, etc.
The assessment was based on a comparative economic evaluation of the
various processes. Some of the factors considered in the study were capital
and operating costs, by-product utilization and/or disposal including 'related
credits or debits, and a pragmatic consideration of the effect or impact of
high-volume production of by-products upon future markets.
As the result of the Phase I effort several candidate processes were
selected for further consideration in Phases II and III. However, the work be-
yond Phase I was not considered to be limited to the chosen processes, if one
or more promising aqueous scrubbing methods were conceived during the
later stages of the program, they would also be analyzed in due course.
It is emphasized that the results reported herein are based on a compa-
rative economic analysis of aqueous scrubbing systems. As such, one can
interpret the results on a relative basis for the systems considered. However,
a comparison of any of the costs cited for a specific process (e. g. , operating
cost per ton of coal consumed) with analogous costs derived from the work of
other investigators should be done with extreme caution. This is particularly
true if a non-aqueous process is being compared with one or more of the
aqueous processes reported here. One can make such comparisons only after
confirmation that the same xnitial cost considerations are used. The present
study is a conservative one in the sense that the process costs are probably high.
-------
II. SUMMARY
A preliminary assessment of aqueous solution methods used for the
removal of sulfur dioxide from waste gases was conducted. Of the more
than twenty processes that were evaluated on an economic basis, only four
were considered as candidates for continued study under Phases II and III.
These are the (1) Zinc Oxide, (2) Ammonia-Hydrazine Exorption, (3) Cominco
Exorption, and (4) Mitsubishi Lime processes. The Zinc Oxide and Ammonia-
Hydrazine Exorption processes both have relatively low capital investment
and operating costs; both also generate sulfur dioxide as the main product.
The Ammonia-Hydrazine Exorption process can also provide anhydrous
hydrazine to the extent that it can be sold, with the unmarketable portion
reusable in the process. The Cominco Exorption process is similar to the
Ammonia-Hydrazine Exorption process, in that sulfur dioxide is recovered
as such. Its capital and operating costs are somewhat higher than those for
the other three candidate processes. Steam and heat-exchange equipment
costs account in large part for the relatively high cost of this process. The
Cominco Exorption process is the only one of the ammonia processes in which
sulfur dioxide is the principal product. The Mitsubishi Lime process has low
capital and operating costs, and produces a high purity gypsum as the major
product. It is the only candidate process that employs a slurry as the ab-
sorbent, although this is not considered a significant process disadvantage.
Two processes were considered as marginal candidates. These are the
Magnesium Oxide and the Manganese Oxide processes. In both, sulfur dioxide
is recovered by calcination of the metal sulfite or sulfate. These processes
were eliminated primarily because they compare poorly with the Zinc Oxide
process, in which a calcination is also involved. In the latter process the
calcination is conducted at a relatively low temperature and yields essen-
tially pure sulfur dioxide; whereas, in the Magnesium and Manganese Oxide
processes, a mixture of products is produced at elevated temperatures.
Both of the latter processes, as they are described in the literature, also
exhibit higher costs than the Zinc 'Oxide Method.
-------
Relatively high investment and operating costs accounted, in large part,
for the elimination of the Haemsch-Schroeder, Ozone-Manganese Ion, Basic
Aluminum Sulfate, Sulfidine, and Wet Thiogen processes. The Mitsubishi
Red Mud process has a low capital cost but a high operating cost; it too was
not considered attractive.
A number of ammonia-based processes were eliminated because their
attractiveness is dependent upon an unrealistically high credit for ammonium
sulfate. These are the Fulham-Simon-Carves, Showa-Denko, Cominco,
Ammonia-Hydrazine, Mitsubishi Manganese Oxyhydroxide, Mitsubishi
Ammoniacal Liquor, and Magnesium Hydroxide processes.
Several processes were not considered attractive because sulfur is not
recovered in any form. These include the Battersea, Howden-I. C. I. (Cyclic
Lime), Mitsubishi Red Mud, and Simplified Lime processes. Another negative
feature of these methods is that the by-products formed are solid wastes which
pose serious disposal or pollution problems.
A number of processes identified during the course of the work were not
evaluated as part of the Phase I effort. Some of these are proprietary, and
the limited data available are not sufficient to permit a satisfactory analysis.
These include the processes under current development by Wisconsin Electric
Power/Universal Oil Products, Combustion Engineering, Bechtel, Ionics/Stone
and Webster, and Wellman Lord (Beckwell). The Kanagawa process uses
water as a scrubbing medium. Very little data were found on the process.
For this reason, and since it is similar to other processes using large
volumes of water as the scrubbing medium, it was not evaluated. The
Guggenheim process was not considered because it is similar to the Cominco
Exorption process except for the method of SO2 recovery which is discussed
later in the text.
The results of the extensive literature review, that yielded nearly 700
papers, reports, patents, etc. , provided much of the information needed for
the evaluation. Typical information included that pertaining to fundamental
absorption phenomena, physico-chemical data, process development and
application, by-product recovery, absorbent regeneration, materials of
-------
construction, process limitations, and economics. Some of the more important
references are cited in this part of the report.
Chemical equilibria calculations were conducted to determine the potential
of the various aqueous scrubbing processes for absorbing sulfur dioxide from
flue gas. The justification for this work was that the results might have served
as a means of readily eliminating processes which theoretically could not be
used for reducing the sulfur dioxide content of the flue gas from 0. 3% to 150 ppm.
Process assessment or analysis would then have been conducted only for the
remaining, more promising systems. However, it was subsequently shown that
any aqueous scrubbing process can theoretically yield an effluent gas which
contains 150 ppm of sulfur dioxide. Accordingly, the only feasible approach to
process selection was that involving a direct, comparative economic analysis.
in. TECHNICAL DISCUSSION
A. LITERATURE REVIEW
An extensive literature survey was conducted to collect available
data related to processes for removing sulfur dioxide from flue gases by aqueous
scrubbing methods. This effort has resulted in the acquisition of an extensive
amount of information; for example, some 500 technical papers alone have been
acquired.
The literature information has been extensively used in carrying out
Phase I of the program. Table 1 lists the processes which have been identified.
Some pertinent information for most of the individual processes is also included
in the table. The majority of the processes listed in the table have been con-
sidered in the economic and technical analysis covered under Phase I; this
aspect of the program will be discussed in detail in subsequent sections of the
report.
The following method of referencing the literature has been used
in this report. In the text itself, references are identified in relation to specific
pertinent material which is presented; the references are identified in the
Bibliography, Part Six. In addition, a separate group of references for
-------
TABLE 1
AQUEOUS SOLUTION SORPTION PROCESSES FOR REMOVING SO2 FROM WASTE GASES
Process
1. Fulham-Simon-Carve 3
Absorber
Makeup
NH,
End Products
(NH4)2S04, S
2. Showa-Denko Ammoniacal NH,
00
3. Cominco
4. Cominco Exorption
Zinc Oxide
NH,
NH,
(NH4)2S04
S02.
SO., CaSO, (im-
2 pare) *
6,7. Howden-I. C.I. (Cyclic Suspension of CaSO-/CaSO4
Lame)/Mitsubishi CaO. Ca(OH)2> (impure)
Simplified Lime or CaCO,
Brief Description
H.SOj added to scrubber effluent
siBesfream and mixture auto-
claved to produce (NH4)2SO4 and
S.
Air passed through scrubber
effluent sidestream to oxidise
(NH4)2S03 to (NH4)2S04.
SO-y evolved from scrubber
effluent sidestream on acidifi-
cation with H,SO. (NH.),SO.
also produced. *t*
SO, stripped from scrubber
effluent sidestream by heat
(steam). (NH4),SO, solution
returned to scrubber.
Addition of ZnO to scrubber
effluent sidestream yields
ZnSO, and Na,SO, solution.
ZnSO," calcine% to3 give SO, and
ZnO, the latter being reused.
solution recycled.
Mixture of CaSO,/CaSO4 sludge
separated from sidestream of
scrubber effluent for disposal.
Original Stage of Development
Utility and Present Status
Flue gas Studied extensively on pilot-
plant scale in England. Not
in current use
Flue gas A 25 -oar test unit in operation
since 19bb at an oil-fired
process-stream plant in Japan.
Smelter In use by Cominco at Trail.
gas B. C. Olm-Mathieson has
used process in treatment of
acid plant tail gases at
Pasadena, Texas plant.
Smelter Used by Cominco at Trail,
gas B. C. during World War IL
Flue gas Pilot plant in U.S. in 1940's.
Prototype in Germany at
about same time.
Flue gas First commercial plant
established in Wales in 1930's.
Second plant built shortly
thereafter in London.
Process currently under study
by Mitsubishi Heavy Industries,
-------
TABLE 1 (Cont'd)
Process
Absorber
Makeup
End Products
Brief Description
Original
Utility
8. Battersea
Thames River CaSOj slurry
water containing
CaO or CaCOj
9. Magnesium Hydroxide Suspension of (NH.),SO.
Mg(OH)2 * 2 4
10. Magnesium Oxide
11. Manganese Oxide
Suspension of
MgO
Suspension of
MnO,
SO,, MgSO.
SO,
1Z. Haenisch-Schroeder Water SO-, CaSO.
(impure)
Scrubber effluent is treated with Flue gas
MnSO, to promote oxidation of
CaSO, to CaSO4 which is more
suitable for river disposal
Liquid portion of stream returned
to scrubber.
Insoluble MgSO. formed from ab- Flue gas
sorpUon is oxidized to soluble
MgSO,.
The Utter is reacted with NH, to
regenerate Mg(OH)2 which is re-
turned to scrubber, (NH4)2SO4
also formed in this step.
Mg(HSO-), formed in scrubber is Flue gas
neutralize? with MgO to form
insoluble MgSO,.
The solid is separated and cal-
cined to regenerate the MgO and
yield SOr
Soluble MnSO. and MnS,O, Flue gaa
formed in absorption process.
Sidestream from scrubber
effluent is heated in autoclave
to precipitate these salts which
are then calcined to form SO,
and MnO,. £
MnO2 reused.
Lome added to scrubber effluent Fine gas
to separate sulfate, formed on
oxidation, as CaSO^.
Liquid fraction is treated to
release SO2>
Stage of Development
and Present Status
British have used process in
several power plants.
Process covered by 1952
British patent. No indication
that process was piloted.
Process used on a pilot plant
scale in USSR prior to 1956.
No information on current
status there. NAPCA is
developing this process under
separate contracts.
Pilot-plant studies conducted
by TVA. Not in current use.
Process developed on small
scale in late 1800'a.
-------
TABLE 1 (Cont'd)
Process
13. Wet Thiogen
Absorber
Makeup
Water
End Products
14. Ozone-Ma Ion
15. Sulfidine
16 Basic Aluminum
Sulfate
Water containing H.SO, (dilute)
O3 and MnSO4 ' *
1/1 Mixture of SO,. Na,SO.
xylidine and water
A1(OH)SO,
SO-. CaSO.
(impure)
17. Ammoma-Hydrazme N2H4
(NH4)2S04
Brief Description
Scrubber efauent treated with
BaS to form S and mixture of
insoluble barium salts.
Clarified liquor returned to
scrubber.
Solids fraction heated to distil
off sulfur; residue heated
further in furnace with carbon
to reduce sulfur salts to sulfide
which is reused.
Ozone added to flue gas promotes
oxidation of SO,.
MnSO. serve* £s catalyst for
oxidation.
Sidestream from scrubber
effluent, containing xylidine
snlfite, heated to release SO..
Xylidine regenerated and re -
turned to scrubber.
Scrubber effluent, containing
A1(OSO,H)SO.., heated to le-
lease SO, ana regenerate
A1(OH)SO-.
Some oxidation to sulfate occurs
in scrubber, CaCO, added to
precipitate CaSO* which is
separated and discarded.
Air passed through scrubber
effluent Sidestream to oxidize
(N rU),S03 to (N,H-),SO .
N_M.'produced in p/oEess
marketed and/or returned to
scrubber system as circum-
stances permit.
Original
Utility
Smelter
gas
Stage of Development
and Present Status
Process abandoned after some
pilot-plant operations.
This work appears to have
been conducted in early 1900's.
Flue gas
Smelter
gas
Smelter
gas
Laboratory work conducted at
TVA laboratories. Not in
current use.
Used in Germany prior and
during World War II period
Not in current use.
Firat commercial plant at
Imatra, Finland, operated
from 1936-41.
Plant at Manchester,
in operation in 1958.
-------
TABLE 1 (Cont'd)
Process
18. Ammonia -Hydrazine
Exorption
Absorber
Makeup
N2H4
19. Mitsubishi Ammomacal NH.
20. Mitsubishi Manganese Suspension of
Oxyhydroxide MnOH(O)
t\>
21. Mitsubishi Lime
22. Mitsubishi Red Mud
Suspension of
CaO or CaCO,
Suspension of
residue from
extraction of
alumina from
bauxite.
End Products
SO.,, (NH4)2S04
(NH4)2S04
(NH4)2S04
CaSO4 (pure)
Mud containing
Na2S04
Brief Description
SO, stripped from scrubber
effluent sidestream by heat
(steam).
Resulting (N-H.l-SO, solution
returned to scruober.
N2H4 produced in process
marketed and/or returned to
scrubber system as circum-
stances permit.
Air passed through scrubber
effluent sidestream to oxidize
to (NH4)2S04.
MnSO. formed in absorption
process treated with NH- and
O,, regenerating the MntOH)O
and producing {NH4),SO4.
CaSO. in scrubber effluent
ooadizfed to CaSO*. Fly ash
removed upstream of SO,
scrubber.
Sodiom alur
i silicate
present in the red mud reacts
with SO, to form Na-SO..
Scrubber effluent discaraed.
Only equipment required is
scrubber.
Original Stage of Development
Utility and Present Status
(Flue gas) Paper study
Flue gas
Flue gas
Flue gas
Flue gas
Processes under development
by Mitsubishi Heavy Industries.
Ltd., Japan. Some pilot-plant
-------
TABLE 1 (Cont'd)
to
Process
23. Wiacon Electric
Power /Uiuve rsal
Oil Products
24. Combustion Engineering
Absorber
Makeup
N.2C03
Water
End Products
Na2S03/Na?S04
CaSO3/CaS04
25. Kanagawa
26. Wellman-Lord
(Bflckwell)
Sea water, natural
underground water,
or alkalized water
KjCO,
27. Bechtel Limestone
28. Ionics /Stone fc Webster NaOH H2SO4(conc.)
Brief Description
Two-stage system employs
direct limestone injection into
furnace in combination with
wet scrubbing utilising a
Na-CO3 solution.
Dolomitic limestone injected
into furnace.
SO, removed by reaction with
calcined limestone.
SO- and fly ash removed in
wafer scrubber.
Process reportedly offers an
improved method for attaining
intimate contact between water
and gas stream.
Process believed to involve
use of aqveous K-SO, as
abeorber. with
Original
Utility
Stage of Development
and Present Status
of isolated
Process involves absorption
by aqueous NaOH, and subse-
quent electrolysis to yield
Flue gas
Flue gas
Pilot studies of individual
systems carried out.
Two-stage system under
develc
Flue gas
Process tested at a mid-
western utility power plant.
C. E. reportedly installing
a $1 million installation at
Union Electric Co. '• coal-
fired electric power plant in
St. Louis County, lOssouri.
Pilot-plant studies (10.6OO cih)
conducted in Japan.
Flue gaa Process under development.
Flue gas Process under development.
-------
TABLE 1 (Cont'd)
Process
29. Guggenheim
Absorber
Makeup
NH,
End Products
Brief Description
30. Diethylene Triamine
(other amines, e. g.,
tnethylene tetramine
and tetraethylene
pentamine have also
received some initial
evaluation)
Diethylene
Triamine
S, CaSO. (impure) SO. stripped from scrubber
efffiient side stream by heat
(steam).
SO- reduced to S by heating
Witt coke.
Product (NH4)2SO4 treated
with lime to regenerate NH,
which is reused, by-product
discarded.
SO,
Amines form solid salts with
SO, which are soluble in water.
SO liberated by heat.
Original Stage of Development
Utility and Present Status
Smelter Pilot plant at Garfield, Utah
gas in 1940's (ASARCO).
Process not in current use.
-------
each of the methods examined under Phase I is provided in this part of the
report (Section III. D.). A concise statement of the contents of each of the
references shown in these groups is included. The groups of references
are not considered to be complete as far as literature coverage is concerned.
The objective was to provide the reader with select information on some of
the more important papers, patents, text books, etc.
B. THEORETICAL CHEMICAL EQUILIBRIA CONSIDERATIONS*
1. Conclusions
Chemical equilibria calculations were conducted to deter-
mine the potential of the various aqueous scrubbing processes for absorbing
SO. from flue gas. It was considered that the results of this work might
have served as a means of readily eliminating processes which theoretically
could not possibly reduce the SO2 content of the flue gas from 0. 3%
(3000 ppm) to 150 ppm. Economic analyses would then have been carried
out for the remaining processes.
The results of the study show, however, that essentially
all of the known aqueous processes can yield an effluent gas which contains
150 ppm of SO,. Accordingly, the only feasible approach to eliminating the
"poorer" processes appears to be by the direct economic analysis route.
2, Comments on the SO2~Removal Efficiency of
Specific Types of Processes
Most of the aqueous processes utilize a scrubbing medium
which is a solution (as contrasted with a suspension). Examples of such pro-
cesses are the ammonia systems (Fulham-Simon-Carves, Cominco, Showa-
Denko, Guggenheim, and Mitsubishi), sodium hydroxide system (Zinc
Oxide), water systems (Battersea, Kanagawa, and Haenish-Schroeder),
and others. The data shown in Table 2 are representative of the results
generated from a series of calculations. Several of these solutions, e. g.,
the ammonia and sodium hydroxide systems, are of sufficiently high basicity
*A discussion that covers some of the more theoretical aspects of the
calculations, the source of thermodynamic data, etc., is presented as
Appendix A in Volume I of this report.
-------
TABIE 2
Scrubbing Liquor
1 M NH3
6. 5 M NH3
6. 5 M NaOH
H20
1/1 H20/Xyhdine
***
6. 5 M Slurry
CaO (CaCOj)
***
6. 5 M Slurry
MgO (MgCO3)
EQUILIBRIA DATA FOR
Major Chemical Input Gat
Constituent (moles)
S02 1
H20 24.2
NH3
so2 i
H20 24 2
NHj
so2 i
H2O 24. 2
NaOH
SO, 1
H2O 24. 2
SO, 1
H20 24.2
Xyhdine
so2 i
H20 24.2
CaO
so2 i
H2O , 24.2
MgO
VARIOUS SCRUBBING
SOLUTIO
Input Chemicals Stack Lots
(moles) (moles)
.
69.0
0.97
-
19.0
1.2
-
17.5
1.0
-
24, 600
-
70.3
8.4
-
16.4
0.95
-
24.8
1.64
0.05
39.1
0. 0004
0.05
34.1
0. 10
0.05
32.9
-
0.05
40.6
0.05
35.3
0.4
0.05
40.6
-
0.05
35.0
-
Scrubber solution temperature is 50°C
Partial pressure of incoming SO, is 2. 28 mm (0. 3%)
**
****
Partial pressure of effluent SO- is 0. 114 mm (150 ppm)
Moles /liter solution
System contains 6. 5 moles of oxide per liter of water
, i
Solid CaSO, only is removed from effluent
Product Capacity
(moles) (moles Alter HO)
0.95
24,600
0.95
59.2
8.8
0.95
.****
0.95
0.95
14.0
1.64
0.17
1.5
1.5
0 0024
0.42
6.5
3.2
-------
to reduce the SO vapor pressure to very low values. In the extreme case
£t
of neutral water, there is a very low concentration of SO- (0. 0024 M)
associated with an SO, partial pressure of 150 ppm at 50 C. As long as
£
the entering scrub water has a concentration of SO2 that is lower than
this value, water is also capable, from an equilibrium standpoint, of at-
taining the desired effectiveness; very large volumes are required, how-
ever. The xyhdine/water system (Sulfidine process) can also be used for
reducing the SO- content to the desired level; the capacity of this system
£t
is not high, however.
An analysis of the presence of CO2 in the systems dis-
cussed in the foregoing shows that a very low level of CO2 is dissolved.
Although alkaline solutions are commonly added to scrubber systems, the
circulating scrubber solution is usually slightly acidic due to prior absorp-
tion of SO,; the acidic conditions thus account for the low level of CO2
solubility.
A number of the other processes, e. g., the I. C. I. -
Basic Aluminum Sulfate, Magnesium Oxide, Howden-I. C. I., Magnesium
Hydroxide, Mitsubishi Simplified Lime, and the Mitsubishi Lime, use
magnesium or calcium oxide or hydroxide slurries as scrubbing media.
In these systems, the formation of solid carbonates is an important con-
sideration as regards the equilibria with SO,. The systems are, however,
£*
also theoretically capable of reducing the SO, content of flue gas to 150 ppm.
Ct
3. Sample Calculation
In order to clarify the significance of the data in Table 2
and also the general procedure used to generate these data, a sample calcu-
lation for a 6. 5 M ammonia scrubbing solution will now be illustrated. The
scrubbing solution at the top of the scrubber is assumed to be at 50°C and
in equilibrium with the exit gases containing 150 ppm of SO_. The 150 ppm
value corresponds to a partial pressure of 0. 114 mm. A calculation of the
equilibria involved shows that a solution with a total ammonia content of
6. 5 M will be 4. 27 M in SO2< The partial pressures above this solution
will be 77. 8 mm H2O, 0. 226 mm NH3, and 0. 114 mm SO2. The volume
-------
of entering flue gas that carries 1 mole of SO, into the scrubber at a
partial pressure of 2. 28 mm (0. 3%) will leave the top of the scrubber
carrying 0. 05 moles of SO2 at a partial pressure 0. 114 mm.
Since 0. 114 mm SO- corresponds to 0.05 moles stack
£t i
'loss, and the stack losses are proportional to the partial pressures, i
77. 8 mm t^O and 0. 226 mm NH. correspond to a stack loss of 34. 1
moles H2O and 0. 10 moles NH,, respectively.
The solution at the bottom of the scrubber containing
6. 5 M total ammonia contains 5. 80 M dissolved SO-, in equilibrium with
0. 3% SCX, in the incoming flue gas. The composition of the solution at the
bottom is 55. 55 moles H2O, 6'. 5 moles NH,, and 5. 80 moles SOg. There-
fore, for each amount of solution containing 0. 95 moles SO, withdrawn
from the scrubber, there is withdrawn 9. 1 moles H,O (55. 6x0. 95/5. 8)
and 1. 1 moles of ammonia (6. 5x0. 95/5. 8).
The capacity of the scrubbing solution is determined
from the difference in SO, concentrations at the top and bottom of the
£•
scrubber. The scrubbing solution contains 5.80 moles SO^/liter H-O at
the bottom of the scrubber and 4. 27 moles SO,/liter H,O at the top. The
difference of 1. 5 moles represents the amount of SO, discharged from
the system for each liter of water in the scrubbing solution.
A mass balance shows the quantity of chemicals that
must be added for each mole of SO, scrubbed. There are 24. 2 moles of
water in the incoming gas stream (7. 25%) while 34. 1 moles are lost in the
stack and 9. 1 moles are withdrawn from the bottom of the scrubber. The
difference, 19. 0 moles, must be added. There must be an addition of
1. 2 moles of NH, to make up for the 0. 10 mole lost in the stack and the
1. 1 moles withdrawn.
-------
C. ECONOMIC ANALYSIS
1. Basis for Analysis
a. General Considerations
i *
The prelimimiry economic analyses reported
herein of processes for removing sulfur dioxide from flue gases of power
plants are based on the following general considerations.
A power plant of 120, 000 kw was selected from a
621#
previous cost study. The cost estimates in the referenced study cover
the Fulham-Simon-Carves, the Zinc Oxide, and the Howden-I. C. I. (Cyclic
Lime) processes, all of which were evaluated on this program. The earlier
work thus provided a sound basis for the present study. The mechanics used
in utilizing these data in the present study are illustrated in the detailed
economic analysis of the Fulham- Simon -Carves Process (see Section III. C.
2. a.). This detail is not repeated in the sections covering the other processes.
Conditions characteristic of the selected power plant are as follows:
• Flue gas contains 0. 3 vol-% SO,
• Removal efficiency is 95% (purified
gas would contain 150 ppm SO,)
• 475, 000 tons of coal are consumed
per year
• Coal requirement is 1 pound per
kilowatt-hour
• Sixty tons 0f coal are consumed per
hour evolving 20 million CFH flue gas
at standard conditions of 0 C and
760 mm Hg
• Operations are based on a 330-day
operating year.
Chemical-reaction yields were assumed to be equi-
valent to 100% conversion unless specifically indicated otherwise. No allow-
ances were made for m-plant or other losses, such as exit losses in the
purified flue gas.
* See bibliography, Part Six, for references.
-------
Chemical consumption and by-product or waste
generation, if applicable, were calculated for the treatment of 1 million
SCF of flue gas. These quantities can be easily factored to a power plant
of any size.
Flue-gas composition is in accordance with the
specifications in the program work statement. This information is repro-
duced below. The quantity of each component contained in 1 million SCF
of flue gas is also given.
Quantity Per Million SCF
Component
N2
C°2
H2°
°2
so2
NO (As NO)
X
Fly Ash
Vol-%
74.9
14.7
7.25
2.8
0.3
0.05
Lb Moles
2086
410
202
78
8.36
1.39
Lb
58, 408
18, 040
3,636
2,496
535
42
83, 157
167
83, 324
wt-%
70.10
21.65
4. 36
3.00
0.64
0.05
99.80
0.20
100.00
At a 95% SO2 removal level, the quantity of SO,
absorbed per million SCF of flue gas is 7. 94 Ib moles or 508 Ibs.
For purposes of simplification, the following as-
sumptions were made in carrying out the analyses:
• All gaseous components other than water vapor
and SO, will pass through the absorber(s) un-
changed.
• The water vapor content of the gas will change.
The exit gas will be saturated with H,O at the
-------
• The untreated flue gas will be available at
300°F, and its pressure will have to be in-
creased to overcome the pressure drop in
the absorption system.
• * The effect of fly ash has not been considered.
• The effect of any required gas cooling on
plume buoyance and dispersal and the need
for any plume reheat have not been considered.
b. Capital Costs
Costs of preparing various types ol capital-cost
estimates are discussed by Bauman. A tabulation from that source
follows:
Costs of Preparing Indicated Estimated
Types
Estimate Type Median Cost ($) Median Cost ($)
(Up to $1 Million) (Over $1 Minion)
Order-of-Magnitude 1000 4000
Study 2000 5000
Preliminary 7000 16000
Definitive 12000 35000
Only order-of-magnitude estimates were utilized
in Phase I. More specifically, fixed-capital type estimates were utilized'
in which the cost of major purchased process equipment was calculated or
determined from established or published data. The additions necessary
to arrive at the values of the direct-plant and fixed-capital costs were de-
termined as percentages of the physical-plant cost and the direct-plant
cost, respectively. In some cases it was adequate to employ Lang's factors
for fixed capital, a procedure which is commonly used in the process indus-
tries for order-of-magnitude estimating purposes?68
Although these estimating procedures are not con-
sidered to be more than + 30% accurate, they serve adequately since the
estimates are used only for obtaining relative costs to determine the feasi-
bility of utilizing various aqueous scrubbing techniques. The major scrubbing
-------
equipment is quite similar in most of the processes. The moat important
variance in equipment requirements is associated with absorbent regenera-
tion, recovery of sulfur in a salable form (S, SO2, H2SO4), and in the
supporting chemical-plant complex - if required.
The factors for fixed capital for three general
processes, as derived by Lang, are as follows:
Thus,
I,
Process
Solid
Solid-fluid
Fluid
EL
Factor
3.10
3.63
4. 74
where: I,, = fixed-capital investment
i
E = purchased equipment cost
L = Lang's factor
The 4. 74 factor for fluid processes was applied in Phase I of this work.
Justification for this selection follows.
These estimate types, as well as many others,
A A.1
appear frequently in the cost literature. Aries and Newton has been
a very popular source of cost information for many chemical engineers.
The method in the referenced publication, described as Fixed-Capital
Estimate, Method 2, was adapted for use on this project. A breakdown
3
of the capital cost items considered by Aries and Newton follows.
-------
Item Fact01:.
Purchased equipment
Installation labor Ot 2
Foundations fe platforms 0> l8
Piping °-86
Instrumentation °'05 to °- 30
Insulation °- °8
Electrical °- 10 tO °'15
Buildings °'30
J^and & yard improvements 0- 10 to °' 15
Utilities °-40
Physical-Plant Cost 3.32 to 3.67
Engineering & Construction (20% of
physical-plant cost) 0.66 0.73
Direct-Plant Cost ' 3.98 to 4.40
Contractor's Fee (7% dir. pit. cost) 0.28 0.30
Contingency (15% dir. pit. cost) 0.60 0.66
Fixed Capital 4.86 to 5.36
It should be noted that these factors are higher than Lang's (4. 74). Even greater
variances are found in the cost literature.
Table 3 is presented to illustrate these variances.
186
Column I is derived from the literature and shows average costs as a ratio
of purchased equipment for large fluid-processing installations. The over-all
fixed capital factor is practically the same as Lang's.
The values shown in columns II, III, and IV were obtained
621
from capital-cost data developed by the Bureau of Mines, all based on plants
with a flue-gas volume of 20 million SCFH containing 0. 30 vol-% SO-, and
assuming 90% removal of the SO-. These also are factored estimates, and
£~ A *t
reference is made in the paper to data from Aries and Newton and others. The
fixed capital factors are substantially lower for all three processes than for
Lang's. This observation is not critical, however, since they compare well with
each other.
-------
TABLE
FACTORED FIXED CAPITAL ESTIMATES
ITEM
I
Haselbarth
& Berk
Large Fluid
Process
II
Nonregenerative
Limestone
Process
III
IV
Field, et al.
Ammomacal
Liquor
Process
Regenerative
Sodium Sulfite
Process
V
Modified
Lang
Fluid
Process
00
Purchased Equipment
Erection Labor
-Foundations JU Platforms
Piping
Instruments
Insulation
Electrical
Buildings
Land & Yard Improvement
Utilities
Receiving, Shipping & Storage
Physical-Plant Costs
Engineering & Construction
Direct Plant Cost
Contractor's Fee
Contingency
Fixed Capital Costs
1.00
( 0.43
0.76
0. 19
-
0. 10
0.23
0. 14
0.57
ge 0. 24
3.66
0.48
4. 14
In Eng. & Const.
0.62
4.76
j 1.43
0. 72
0.05
-
0.05
0.30
0.13
0.11
Misc. 0.11
2.90
0.75
3.65
-
0.35
4.00
| 1.43
0.67
0.05
-
0.05
0.30
0. 13
0. 10
Misc. 0. 11
2.84
0.74
3.58
-
0.34
3.92
| 1.43
0.60
0.05
0.01
0.05
0.30
0. 10
0. 11
Misc.O. 11
2.76
0.72
3.48
-
0.33
3.81
1.00
0.25
0. 18
0.76
0. 15
0.08
0. 10
0.25
0.13
0.40
-
3.30
0.66
3.96
0.19
0.59
-------
Many more variances in fixed capital factors (as
a percent of purchased equipment cost) can be cited from the literature.
The examples given, however, should suffice to indicate the importance
of having common bases whenever order-of-magnitude estimates are to
be compared with each other. The 4. 74 fixed capital factor was selected
for Phase I of this project. Volumn V of Table 3 indicates the individual
values used.
Working capital usually does not exceed 10-15%
of fixed capital. In the Phase I effort of this program, the equivalent of
10% of fixed capital was assigned as working capital.
The sum of the fixed capital cost and the working
capital provided the total investment used in the cost estimates of Phase I.
The capital cost in terms of dollars per kilowatt generating capacity was
considered for each process evaluation.
It is important to remember that the capital in-
vestments obtained according to this factored order-of-magnitude method
should noc be construed as absolute values. They are to be used only to
indicate the relative investment costs of the processes under consideration
in Phase I. Comparison of these capital investments with others appearing
in the literature should be made with extreme care due to the large varia-
tion in the magnitude of factors being used in order-of-magnitude cost
estimates.
The capital investments in the Bureau of Mines
costs were adjusted using the Chemical Engineers' Plant Cost Index.13 In
a typical case, the cost of the major items installed (from the Bureau of
Mines report) was divided by 1. 43 to obtain the purchased items coat.
This value was updated from 1957 to 1967 by increasing the cost by 12.4%.
Preliminary capital-cost estimates made in
Phase III on candidate processes were done in a more accurate manner
in accordance with the requirements set forth in the program work state-
ment.
-------
c. Operating Costs
The operating costs include all of the direct,
indirect, and fixed charges which generally occur in manufacturing
operations. The selection of these cost elements is straightforward.
However, assignment of unit values to the cost elements is done on a
uniform basis in all of the ope rat ing-cost estimates. In other words,
labor rates, cost of power per kwh, etc., are the same in all estimates
in order to simplify the task of evaluating the relative economics of the
various processes. A wide variation of values has been assigned by dif-
ferent workers to these cost elements; some of these were determined by
actual costs and others were estimated on the basis of past experience.
j
For the purpose of this program, it is only necessary to select realistic
values for a hypothe ical situation or location. A brief treatment of these
cost elements follows.
(1) Direct Costs
(a) Raw Materials, Processing
Chemicals, and Catalysts
The costs of these materials have
been obtained from standard current sources*, such as the Oil, Paint and
Drug Reporter, (OPDR) or from direct quotations. The initial purchase
of catalysts, if any, should be capitalized and listed in the plant invest-
ment. Only replacement purchases are considered as part of the operating
cost. The raw material prices used in the operating cost estimates are
shown in Table 4.
(b) Direct Labor
The annual wage of hourly workers is
based on 2000 hours. Actually, slightly less than 2000 hours per man-year
are charged to direct labor, with holidays, etc., charged to payroll burden.
Wage rates vary with the industry and locality. A rate of $3. 00 per hour is
used in these estimates.
*The use of the term "current prices" throughout this report refer
to OPDR prices as of February, 1968.
-------
TABLE 4
RAW MATERIAL PRICES
All prices obtained from the Oil, Paint and Drug Reporter,
February 5, 1968, unless otherwise noted.
Raw Material
Ammonia, anhydrous
Ammonium sulfate
Calcium sulfate
Calcium sulfite
Chlorine
Coke
Hydrazine, anhydrous
Lime '
Limestone, 100 mesh
Magnesium oxide, 95%
Magnesium sulfate
Manganese sulfate, 75%
Oleum, 20%
Ozone
Soda ash
Sodium sulfate
Sulfur
Sulfur dioxide, liquid
Sulfuric acid, 100%
Xylidine
Zinc oxide
Price, Dollars
60 per ton
31 per ton
Oa
oa
0. 0325 per Ib
19 per ton
2.95 per lbc
15. 50 per ton
10 per ton
61 per ton
2.45 per 100 Ibs
100 per ton
34. 90 per ton
0. 105 per lbe
1. 60 per 100 Ibs
oa"
38 per long ton
0. 0345 per Ib
33.40 per ton
0.45 per lbf
0. 1525 per Ib
a. Waste products assumed"at no value.
b. Bibliography reference 653.
Olin Mathieson.
United Lime Division, Flintkote Co.
e. Bibliography reference 424.
f. DuPont (mixed xylidines).
c.
d.
-------
(c) Operating Supervision
i
With first-line supervision receiving
approximately 25% more than direct labor, an annual rate (salary) of
$7, 800 is assigned for this cost element. The plant or area superintendent
rate is $12, 000 per year. The quantity of manpower required, both direct
labor and supervision, is estimated for each process under consideration.
(d) Maintenance
Maintenance materials, labor, and
supervision are grouped in this cost element. Due to lack of adequate data,
the total annual cost of maintenance is set at 5% of the fixed capital invest-
ment.
(e) Plant Supplies
Plant supplies are equivalent to 15% of
maintenance cost.
(f) Utilities
Utility unit costs are extremely variable
in practice and are usually controlled by the plant location and size. The
following unit costs were used:
Utility Cost in Dollars
Steam, M Ib 0. 50
Heat credits & debits, MM Btu 0. 50
Power, kwh 0. 006
Raw water, M gal 0. 10
Recirculated water, M gal 0. 05
Fuel oil, gallon , 0. 10
(2) Indirect Costs
(a) Payroll Burden
This account covers social security,
workmen's compensation, vacations and holidays, contributions to pensions,
-------
life insurance and profit-sharing, etc. Consequently, the cost of these
fringe benefits is rising- steadily and is usually between 15 and 25% of
base labor costs. In this work, this cost was shown as 20% of total base
labor.
(b) Plant Overhead
The cost of maintaining service func* '
tions such laboratory, purchasing, warehousing, engineering, etc., was
not individually estimated. Plant overhead was considered to be 50% of
labor, maintenance, and supplies.
(c) Packaging and Shipping
These costs were not considered in
Phase I.
(d) Waste Disposal
57
This cost was set at 20 mills per
ton mile assuming a maximum haulage of 200 miles, or $4 per dry ton.
Transporting waste shorter distances would'not affect this cost signifi-
cantly since handling costs remain constant.
(3) Fixed Costs
(a) Depreciation
The following depreciation information
issued as guide-lines by the U. S. Treasury Department is applicable to
this project:
• Electrical utility (steam producing plant) 28 years
• Chemical plant 11 years
• Factory buildings 45 years
Although the SO2 removal equipment will represent additions to power plants,
it is doubtful that the useful life of this process equipment would be 28 years.
The 11-year period for a chemical plant is probably more realistic for SO,
removal systems. For simplicity, a 10-year straight-line depreciation which
ignores the 45-year write-off on buildings was applied to the fixed capital.
-------
(b) Taxes
A value of 2% of the fixed capital in-
vestment has been used in this project.
(c) Insurance
Although a small item, 1% of fixed
capital investment was employed to cover plant insurance costs.
An operating cost estimate summary was com-
pleted for each of the processes evaluated. These are presented in the
respective sections in which the processes are considered.
(d) Cost of Capital
Although the interest cost associated
with borrowing capital can be an appreciable operating cost item (typically
8% interest in today's market), this cost item was not included in deriving
process operating costs in this study, so that said derived costs are some-
what low due to this exclusion.
d. Profitability
The profitability of each process was considered.
A simplified approach was applied for this part of the analysis. The com-
monly used indexes, such as payout time, return on investment, etc.,
were not applied. The data are presented simply as a loss or profit in
terms of dollars per ton of coal burned and mills per kwh generated at
various levels of by-product sales (when applicable).
The magnitude of by-product credits applied to the
operating costs can severely affect the profitability of a process. In some
cases, the by-products would have a definite impact on the market and on
the economy of the United States.
The various by-products produced in these systems
are ammonium sulfate, calcium sulfite, calcium sulfate, sulfur, dilute sul-
furic acid, magnesium sulfate, sulfur dioxide, sodium sulfate, and anhydrous
hydrazme. Conversion of even a small part of the SO^ emissions from
fossil-fuel, power-generating plants to any of these by-products would pro-
vide huge quantities of materials.
-------
Several of these by-productB have substantial
value in today's market. These include: ammonium sulfate, sulfuric acid,
sulfur, sulfur dioxide, and anhydrous hydrazine. However, their ultimate
value could be, seriously affected by the large tonnages which would be pro-
duced by the removal of SO2 from flue gases. The rest of the by-products
listed were considered to be waste materials with no value; in addition, a
cost is associated with their disposal.
A discussion of the market considerations of the
by-products follows.
(1) Ammonium Sulfate
Ammonium sulfate was the major nitrogen
fertilizer in the United States from 1923 to 1947. Most of it was produced
from coke oven by-product ammonia and sulfuric acid. Since that time it
has been replaced in its No. 1 position by ammonium nitrate and more re-
cently by anhydrous ammonia. The relatively low nutrient content of
ammonium sulfate has been the major reason for its reduced acceptance
in the fertilizer market in the U. S. A. However, it is still the leading
nitrogen fertilizer in world-wide areas of relatively low agricultural
development.
Production capacity of anhydrous ammonia
increased at a steady rate in the late 1950's; however, the most dramatic
increase in capacity has occurred since 1964. The capacity of synthetic
ammonia plants in the U. S. A. in 1964 was 7, 838 M short tons with a pro-
jected increase to 17, 246 M short tons by January 1968. This great expan-
sion was due to the technological improvements in equipment, resulting in
the construction of very large ammonia plants; the effect has been a very
significant reduction in production costs. Over 75% of the ammonia pro-
duced in the United States today is used in fertilizer. It is used not only
as a fertilizer intermediate but is now the leading direct-application
nitrogen fertilizer in the United States, due mainly to its low cost.
-------
The total production of ammonium sulfate
in the United States in 1963 was 1, 823, 000 tons. A 120-megawatt power
plant would generate enough SO2 in its flue gas to produce 83, 000 tons
per year of ammonium sulfate; a 1400 megawatt station, the plant size
designated for Cases I and II in Phase III of this study, would produce
622, 500 tons per year of ammonium sulfate, or one -third of the annual
Q. S. production.
It is evident that production of ammonium
sulfate from the SCX, available in the flue gas of several 1400 megawatt
power plants would severely upset the ammonium sulfate market. Thus,
this by-product cannot be considered as a desirable one based on present
conditions, especially when it is the only or the major by-product in a
process.
Perhaps a reasonable assumption is that
ammonium sulfate could capture a part of the fertilizer market now served
by anhydrous ammonia if their price relationship improved. The nutrient
content (nitrogen) of ammonia is 82. 3% whereas ammonium sulfate contains
21. 2%. One unit of nitrogen, which is equivalent to 20 Ib of nutrient, costs
twice as much in ammonium sulfate as in anhydrous ammonia at today's
prices of $31 per ton for ammonium sulfate and $60 per ton for ammonia:
In NH3 In
Units of nutrient, per ton 82.3 21.2
Fertilizer cost, $ per ton 60 31
Cost per unit, $ 0.73 1.46
Thus, to compete with ammonia on the basis of nutrient content, ammonium
sulfate would have to sell for $15. 50 per ton when the price for anhydrous
ammonia is $60 per ton. On this basis, any reduction in anhydrous ammonia
price would mean a decrease in the ammonium sulfate price. These price
relationships are also affected by geographical location and on specific situa-
tions. The total net cost of the nutrient applied to the farmer's soil is
probably the determining factor affecting the market price for which am-
monium sulfate would have to be sold to be competitive with anhydrous
ammonia.
-------
The world-wide market is also a factor
since ammonium sulfate is still a major fertilizer in Asia and Europe.
Shipping costs, however, would affect the economics. As a matter of
interest, however, a recent newsletter announced the sale of 2. 25
million metric tons of Japanese ammonium sulfate valued at $69. 6 million
to mainland China. This is equivalent to 2. 475 million short tons at
$28. 12 per ton.
It is apparent that assignment of credit to
ammonium sulfate produced as a by-product in a SO2 removal system
must be done with care. It is unlikely that an adequate market would
exist for the large tonnages of ammonium sulfate which could be produced
from this source. In this connection it should be noted that the recovery
of SO, from flue gas in systems which produce ammonium sulfate does
Lt
not provide a compound which is valuable for its sulfur content. Am-
monium sulfate used as fertilizer is valuable primarily for its nitrogen
content; although sulfur is a plant nutrient, its function as a component
of ammonium sulfate as fertilizer is of secondary importance.
(2) Sulfur and Sulfur Dioxide
Over 70% of the sulfuric acid produced in
the United States has been from elemental sulfur. In 1962 the production
capacity as 100% acid was 26, 000 M tons. Actual production as 19, 000 M
tons of which 13, 000 M tons were produced from elemental sulfur.
Sulfur production in 1961 was 7, 200 M long
tons of which 5, 300 M long tons were derived from native sulfur. The
latter quantity is equivalent to more than 18, 000 M tons of sulfuric acid.
The availability of naturally occurring brimstone has been severely re-
duced during recent years. This shortage has resulted in cutbacks in
sulfuric acid production in some cases. Sulfur producers have found it
necessary to ration the sulfur supply during the past year. This situation
has resulted in a rapid rise in the price of crude sulfur from approximately
$25 per long ton in 1965 to the current price of $42 per long ton. It has
also created renewed interest in finding new sources of sulfur for sulfuric
acid production.
-------
The recovered SO, could be liquetied as
shown in the cost estimates; it could also be converted at the power plant
site to sulfuric acid or reduced to elemental sulfur depending on the
economics in the local situation. Conversion to sulfuric acid would pro-
bably provide the most economic route.
(3) Anhydrous Hydrazine
Anhydrous hydrazine is a by-product in
two of the processes included in this study. Since a dilute aqueous solution
of hydrazine is the absorbing medium in these processes, the hydrazine
can either be recycled or withdrawn as a by-product. The quantity of
anhydrous hydrazine removed from the system can be regulated; the
specific quantity would be controlled by the anhydrous hydrazine market
existing at that time.
Hydrazine is marketed as anhydrous hy-
drazine, 64. 0% aqueous hydrazine (hydrazine monohydrate), 54. 5%
aqueous hydrazine (85% hydrazine monohydrate), and 35% hydrazine. It
is also available in the form of various salts.
(4) Calcium Sulfate
Gypsum, calcium sulfate, dihydrate, is
used extensively in the manufacture of wallboard, and thus is directly
dependent on the building industry. Lesser quantities are used in making
plaster, cement, in soil and water conditioning (purification and clarifica-
tion), and in the beer industry for pH control. The bulk rate for 97% pure
gypsum is $4. 25 to $4. 50 per ton, at the mill. Anhydrous calcium sulfate
is valued at $24/ton, in bulk, F. O. B. mill, but the consumption of this
material is relatively small, the chief use involving water clarification.
The present brief discussion will therefore be restricted to a considera-
'tion of the gypsum market.
-------
The following table shows the world production of
656
gypsum during the period 1960-1962:
Production. M Short Tons
North America
South America
Europe
Asia
Africa
Australia/New Caledonia
The gypsum market has been notably stable.
Several of the processes which have been considered
for the removal of SO, from flue gases afford gypsum as a product. However, in
many cases (e.g., in the Howden-1. C. I. process), the product is contaminated
with fly ash and cannot be credited. In the Mitsubishi Lime process a high purity
product is obtained. Thus, in this process the gypsum produced was initially
considered not as a waste material, but as a by-product with value. The Phase III
evaluation, however, showed a dim forecast for marketing pure gypsum in the
United States.
I960
16,445
452
24, 130
3,946
933
651
46, 560
1961
16,135
542
26,135
3,843
982
681
48,320
1962
16,753
584
26, 720
4,291
934
683
49, 965
-------
2. Detailed Economic Analysis
a. Fulham-Simon-Carves Process
(1) Process Description
The Fulham-Simon-Carves process involves the
scrubbing of flue gas with an aqueous solution of ammonium salts, including the
sulfite, bisulfite, and thiosulfate. Gas works liquor (an aqueous solution con-
taining ammonia and hydrogen sulfide) or synthetic ammonia is added to the
circulating scrubber liquor at a rate corresponding to the rate of sulfur dioxide
absorption, and a portion of the liquor is continuously removed for product
recovery. The latter involves a preliminary filtration to remove fly ash,
followed by an autoclaving step in which the clear liquor is heated in the presence
of sulfunc acid to convert bisulfite to sulfate and elemental sulfur. After the
withdrawal of molten sulfur from the autoclave the aqueous ammonium sulfate
solution is concentrated in a vacuum evaporator, and the precipitated crystalline
solid separated by centrifugation, and dried. A flow diagram for the Fulham-
Simon-Carves process is shown in Figure 1.
(2) Process Reactions
The reactions shown below apply to the case where
gas-works liquor which contains hydrogen sulfide is used as the source of
ammonia. When synthetic ammonia is used, the reactions which show the for-
mation of ammonium thiosulfate in the scrubber and the subsequent reaction of
the material with ammonium bisulfite in the autoclave are not applicable. As
will be indicated, only synthetic ammonia is considered in this study.
NH4HS03
3
(NH4)2S03 + 1/2
2H2S + 2 NH4HS03
Autoclave: H2SO4
2 NHHS0 + (NH)S0 - - - 2-» 2 (NH)SO + 2S
HS0
2 NH4HS03 + (NH4)2S03
24
= - * 2 (NHSO + S
-------
Reactant Chemicals Per Million SCF Flue Gas Processed
PURIFIED GAS
FLUE GAS
A Stream
1
2
3
4
5
6
Component
S02
so2
NH3
20% Oleum
S
(NH4)2S04
Lb Moles
8.36
0.42
14.4
0,54
1.30
7.18
IDs.
535
26.8
244 '
5LO
4L6
948
t
HgO
•< NH3
SCRUBBER
FILTER PRESS
AUTOCLAVE VACUUM CENTRIFUGE
EVAPORATOR
DRYER
FULHAM-SIMON-CARVES PROCESS : FLOW DIAGRAM
-------
(3) Availability of Ammonia
For many years ammonia was obtained commer-
cially as a by-product in the manufacture of coke and gas from coal. This was
the situation when the Fulham-Simon-Carves process was developed some
thirty years ago. Although some ammonia is still produced in the U.S. in this
manner, the most important method now is the Haber process which involves
synthesis from nitrogen and hydrogen.
In the United States, most of the by-product
ammonia is produced in coke-oven plants. This ammonia is recovered as
ammonium sulfate, ammonia liquor, and/or ammonium phosphates. In 1961,
51 plants produced sulfate; 10 plants, ammonia liquor; 3 plants, diammonium
phosphate; and 1 plant, monoammonium phosphate.
In 1964, the synthetic ammonia production capacity
was 7, 838 M short tons of ammonia with a projection of a 17, 246 M short-ton
/ e I
capacity in 1968. Production in 1964 was as follows:
M Short Tons of NH3
Synthetic, anhydrous 7,508.5
Synthetic, aqueous 62. 9
By-product, as (NH4)2SO4 175.3
By-product, as liquor 16. 3
By-product, as phosphates 11.8
It is not likely that the by-product ammonia now con-
verted to ammonium sulfate would be available as an absorbent for the removal of
SO- from flue gas. Most of the ammonium sulfate is produced at steel mills and
is used to solve a waste pickle-liquor disposal problem; the ammonia is reacted
with the waste pickle liquor to produce ammonium sulfate. It will be shown in the
following section that the quantity of by-product ammonia liquor available (16. 3
thousand short tons in 1964) would not be adequate for the removal of SO- from
the flue gas generated by one small existing power plant. For this reason, the
Fulham-Simon-Carves process using by-product ammonia as an absorbent was not
considered. The economic evaluation which follows was therefore based on
the use of anhydrous ammonia.
-------
(4) Chemical Requirements and By-Product Yields
621
Field, et al. used pilot-plant data which were
factored for a 20 million SCFH system designed to remove 90% of the SO2 from
flue gas containing 0. 3 vol-% SO2. The quantities of reactants and by-products
were based on the use of anhydrous ammonia and not gas works by-product
ammoniacal liquor. In the calculations shown herein, a purified flue gas con-
taining 150 ppm of SO, will require 95% removal of the SO2< Although ammonium
thiosulfate is not formed when using anhydrous ammonia, the addition of sulfuric
acid to .the autoclave is still required to convert the ammonium sulfite/bisulfite
to ammonium sulfate and sulfur.
Table 5 shows the raw material and by-product
requirements for treating one million SCF of gas. Data for three different sized
plants, viz., 20 million SCFH, 0. 5 million SCFM and 2. 5 million SCFM, are also
shown in the table. The latter two plant sizes were evaluated in Phase III.
These quantities are based on the assumption of full utilization of ammonia, with
no allowance for ammonia loss in the treated flue gas. Ammonia losses reported
in the pilot-plant operations were 3 to 8% of the make-up ammonia. Ammonia
losses of only 0. 3 to 0.4% of make-up have also been reported in a pilot-plant
673
operation.
It is evident that the 16, 300 short tons of by-product
651
ammonia available in 1964 is not sufficient to satisfy the requirements of any of
the plants which were considered.
The theoretical quantities of reactant chemicals in-
volved with the treatment of 1 million SCF gas are also shown on the flow diagram,
Figure 1.
(5) Cost Estimate
(a) Capital Costs
The capital cost estimate of the Fulham-Simon-
Carves process in the Bureau of Mines report contained the following information:
Fixed capital $4,495, 800
Working capital 449, 600
Total Investment: $4, 945, 400
-------
TABLE 5
FULHAM-SIMON-CARVES PROCESS USING ANHYDROUS AMMONIA:
CHEMICAL REQUIREMENTS & BY-PRODUCT YIELDS
Quantity per Million
Tons per Year
aw Materials
SO, (flue gas)
b
NH, (anhyd. )
Oleum, 20%
SCF Flue Gas
Lb Mole Lb
7. 94 508
14. 36 244
0.54 51
20 Million
SCFH Flue Gas
40,230
19,325
4,040
0. 5 Million
SCFM Flue Gas
60, 350
28, 990
6,060
2. 5 Million
SCFM Flue Gas
301,730
144, 940
30, 300
y-Products
(NH4)2S04
7. 18
1. 3
947
42
75,000
3,325
112,500
4,990
562, 500
24,940
-------
The cost of major items installed is $1, 637, 240 of this total. Dividing this value
by 1.42 provides the purchased equipment cost, or $1, 144, 900. In this case, the
fixed capital cost, $4,495, 800, is 3. 92 times the purchased equipment cost. The
fixed capital cost in the present work was taken as 4. 74 times the purchased equip
ment cost. This variance is due to the selection of, and the values applied to,
the specific factors making up the fixed cost.
The purchased equipment cost of $1, 144, 900
in 1957 was corrected to November 1967 costs using the C.E. Plant Index. This
amounted to a 12.4% increase, or a total of $1, 286, 900. This amount was
factored according to established standard procedure, resulting in a total in-
vestment of $6, 704, 700, or $55.87/kw capacity. See Table 6.
(b) Operating Costs
The raw material requirements tabulated in
Table 5 for the Fulham-Simon-Carves process were based on the assumption that
anhydrous ammonia would be used. As mentioned earlier, the supply of by-
product ammonia would not be adequate; therefore, operating cost estimates were
not prepared on the by-product ammonia system. The total operating cost per
year was $3, ?.21, 300 or $6. 78/ton of coal and 3. 39 mill/kwh. See Table 7.
Raw material and chemical costs were based
on the data given in Table 5.
Direct labor was assumed to be the same as in
the Bureau of Mines report, i.e., 5 men per shift x 4 shifts = 20 men. Supervisior
was assumed as one foreman per shift and one area superintendent.
The utility requirements given in the Bureau of
Mines report were factored to adjust for the change from 90% SO2 removal to 95%.
Steam: 364, 000 M Ib per year
Power: 12, 236, 000 kwh
Make-up Water: 200, 000 M gal per year
Circulating Water: 975, 000 M gal per year
In the Bureau of Mines study it was assumed
that the by-products would be handled in bulk. Costs are not included for bags,
labor or bagging equipment. Costs m the present study will similarly be based
on bulk materials.
-------
TABLE 6
FULHAM-SIMON-CARVES PROCESS:
CAPITAL COST ESTIMATE SUMMARY
ITEM FACTOR COST - $
1. Purchased Equipment 1.00 1,286,900
2. Erection Labor 0.25 321, 700
3. Foundation & Platforms 0.18 231,600
4. Piping 0. 76 978,000
5. Instruments 0.15 193.000
6. Insulation 0.08 103,000
7. Electrical 0. 10 128,700
8. Buildings 0.25 321,700
9. Land & Yard Improvements 0.13 167, 300
10. Utilities 0.40 514,800
Physical-Plant Cost 3. 30 4,246,700
12. Engineering & Construction 0.66 849,400
13. Direct Plant Cost 3.96 5,096,100
14. Contractor's Fee 0.19 244, 500
15. Contingency 0. 59 759, 300
16. Fixed Capital Cost 4.74 6,099.900
17. Working Capital, 10% 0.47 604,800
Total Investment 5.21 6,704, 700
18. Capital Requirements
$/kw capacity 55. 87
-------
TABLE 7
FULHAM-SIMON-CARVES PROCESS: OPERATING COST ESTIMATE SUMMARY
Fixed Capital Cost: $6, 099, 900
ITEM
1. Raw Materials & Chemicals
2. Direct Labor
3. Supervision
4. Maintenance, 5% of Fixed Capital
5. Supplies, 15% of Maintenance
<>. Utilities
7. Other
8. TOTAL DIRECT COST
9. Payroll Burden, 20% of 2 & 3
10. Plant Overhead, 50% of 2, 3, 4 & 5
11. Pack & Ship
12. Waste Disposal
13. Other
14. TOTAL INDIRECT COST
i
15. Depreciation, 1Q % Fixed Capital/Yr
16. Taxes, 2% of Fixed Capital
17. Insurance, 1% of Fixed Capital
18. Other
19. TOTAL FIXED COST
20. TOTAL OPERATING COST
21. COST: $/Ton of Coal
22. Mill/kwh
6.78
3.39
TOTAL $
1.300.500
120.000
43.200
305.000
45.800
324.200
2.138.700
32. 600
257.000
289. 600
610.000
122.000
61.000
793.OOP
3.221.300
40.37
3.73
1.34
9.47
1.42
10.06
66. 39
1.01
7.98
8.99
18.94
3.79
1.89
24.62
100.00
23. BY-PRODUCT CREDIT
(SEE FIGURE 2 )
-------
(c) Profitability
Figure 2 indicates the profitability of the
Fulham-Simon-Car /es process on the basis of selling the sulfur produced at
$33. 00 per long ton , with ammonium sulfate offered at variable prices.
On this basis, the estimates indicate that
the cost of removing SO, from the flue gas of a 120, 000 kilowatt plant may be
£*
in the range of $4-5 per ton of coal or 2. 0 -2.5 mills per kilowatt hour. It is
probable that more detailed cost estimates applied to larger power plants would
indicate a reduction of this cost.
With reference to the effect of stack losses on
operating cost and profitability, it is of interest to note that each 1% of makeup
ammonia equivalent lost through the stack wculd increase the operating cost by
$0. 024 per ton of coal burned.
b. Showa-Denko Ammoruacal Process
(1) Process Description
The Showa-Denko Ammoniacal process involves the
scrubbing of flue gas with an aqueous solution of ammonium salts. In this
process, make-up ammonia is injected into the flue gas to reduce corrosion up-
stream of the scrubber, whereas in the Fulham-Simon-Carves process
ammonia is introduced directly into the scrubber. The scrubber itself is
similar to that employed in the Fulham-Simon-Carves process. A portion of the
circulating liquor is continuously removed for product recovery. The off-stream
liquor is filtered to remove fly ash, treated with ammonia to convert any bisulfite
present to sulfite, and then oxidized with air to convert the sulfite to sulfate.
The ammonium sulfate solution is then concentrated in an evaporator-crystallizer,
and the precipitated crystalline solid separated by centrifugation, and dried. A
flow diagram for the Showa-Denko Ammoniacal process is shown in Figure 3.
-------
£•* W
1.84
£ 1-61
1 1.38
«•;
c
n
v_ 1 1C
O) 1. IJ
Q_
«/»
S 0.69
a.
0.46
0.23
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£ 25
120
ol
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ximum price -
jivaient to value
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Basis : • 120 megawatt power plant
•20 million cfh ftue gas
•Sale of 75,000 tons (NH4) S04 and
3326 tons sulfur per yesr
Current Prices
Sulfur
: $31 / ton
: $38 / long ton
2345
Cost - Dollars Per Ton Coal
0.5 1.0 1.5 2.0 2.5 3.0 3.5
Cost - Mills Per Kilowatt Hour
FULHAM-SIMON-CARVES PROCESS USING ANHYDROUS AMMONIA r PROFITABILITY
-------
Reactant Chemicals Per Million SCF Flue Gas Processed
PURIFIED GAS
HjO
FLUE GAS
A Stream
1
2
3
4
Component
S02
S02
NHj
(NHJ S04
HO H
Lb Moles
8.36
a 42
15.9
7.94
Us
535
26.8
270
1048
T
SCRUBBER
SLUME
FILTER MIXER
PRESS
SPENT
AIR
I
OXIDIZER
VACUUM
EVAPORATOR
CENTRIFUGE
SHOWA - DENKO AMMONIACAL PROCESS : aOW DIAGRAM
Figure 3
-------
(2) Process Reactions
Scrubber:
NH3 + H20 + S02 •» NH4HS03
2NH3 + H20 + S02 •• (NH4)2S03
(NH4)2S03 + 1/2
Mixer:
NH4HS0
Oxidize r:
(NH4)2S03 + 1/2 02 * (NH4)2SO4
The available data indicate no pertinent difference
between the absorption system used in the Fulham-Simon-Carves process and
that in the Showa-Denko process when anhydrous ammonia is used. Therefore,
it is reasonable to assume that raw material requirements and by-product
output would be the same. Equipment requirements would vary only in the
oxidation step of the by-product liquor stream. This equipment variance,
which affects capital cost, direct labor, and utility charges, will be indicated
in the cost estimates which follow.
(3) Cost Estimate Detail
Since very little data are available for this procesa
it was assumed that the Fulham-Simon-Carves process plant, with minor modifi-]
cations, could be used for the Showa-Denko Ammoniacal system.
(a) Capital Costs
The major change for this system relative to
the Fulham-Simon-Carves system is the elimination of the autoclave step and the
addition of a facility for air-oxidation of the sulfite sidestream removed for by-
product recovery. The following equipment would not be needed:
Autoclaves - 2
H2SO^ storage tank
Oleum storage tank
Liquor pump to autoclave - 2
-------
Although details were not available concerning the oxidation equipment required,
the following were assumed to be needed:
5000-gal aeration tank
Air compressor
The reduction in purchased equipment was estimated as $81, 300. Table 8
summarizes the capital cost estimate for this process and shows a total
investment of $6, 281, 100, or $52. 34/kw capacity.
(b) Operating Costs
The raw material requirements for the
Showa-Denko process are listed in Table 9. It should be noted that anhydrous
ammonia is the only raw material required and that ammonium sulfate is the
only by-product.
Direct labor is reduced by one man per shift
relative to that required in the Fulham-Simon-Carves process due to elimi-
nation of the autoclave operation. Supervision requires one foreman per shift
and one area superintendent.
Utility requirements are the same as those
needed in the Fulham-Simon-Carves process except for a reduction in steam
cost due to elimination of the autoclave.
Steam: 229, 000 M Ib per year
Power: 12, 236, 000 kwh per year
Make-up Water: 200, 000 M gal per year
Circulating Water: 975, 000 M gal per year
The total operating cost per year was estimated
at $3, 012, 100, or $6. 34/ton of coal and 3. 17 mill/kwh. See Table 10 for details.
(c) Profitability
Figure 4 shows the profitability of this
process, which indicates a cost of $3-4 per ton of coal or 1.2-2.0 nulls per
kilowatt hour.
-------
TABLE 8
SHOWA-DENKO PROCESS: CAPITAL COST ESTIMATE SUMMARY
ITEM FACTOR COST - $
1. Purchased Equipment 1.00 1.205.600
2. Erection Labor 0.25 301.400
3. Foundation fc Platforms 0. 18 217.000
4. Piping 0.76 916.300
5. Instruments 0. 15 180.800
6. Insulation. 0.08 96.400
7. Electrical 0. 10 120.600
8. Buildings 0.25 , 301.400
9. Land fc Yard Improvements 0. 13 156. 700
10. Utilities Q.4Q 482.200
U. Physical-Plant Cost 3.30 3.978.400
12. Engineering fc Construction p. 66
13. Direct Plant Cost 3.96
14. Contractor's Fee • 0. 19 229. 100
15. Contingency 0.59 711.300
16. Fixed Capital Cost 4.74 5.714.500
17. Working Capital, 10% 0.47 566.600
Total Investment 5.21 6.281. 100
18. Capital Requirements
-------
TABLE 9
SHOWA-DENKO AMMONIACAL PROCESS:
CHEMICAL REQUIREMENTS & BY-PRODUCT YIELDS
Quantity per Million Tons per Year
SCF Flue Gas 2Q Mlllion 0 5 i^m^ 2. 5 Million
i,w Materials Lb Mole Lb SCFH Flue Gas SCFM Flue Gas SCFM Flue Gas
LO- (flue gas) 7.94 508 40,230 60,350 301,730
I'
*H, (anhyd.) 15.88 270 21,380 32,070 160,350
-Product
7.94 1048 83,000 124,500 622,500
-------
TABLE 10
SHOWA-DENKO PROCESS: OPERATING COST ESTIMATE SUMMARY
Fixed Capital Cost: $5, 714, 500
ITEM
1. Raw Materials & Chemicals
2. Direct Labor
3. Supervision
4. Maintenance, 5% of Fixed Capital
5. Supplies, 15% of Maintenance
6. Utilities
7. Other
8. TOTAL DIRECT COST
9. Payroll Burden, 20% of 2 & 3
10. Plant Overhead, 50% of 2, 3, 4 b 5
11. Pack & Ship
12. Waste Disposal
13. Other
14. TOTAL INDIRECT COST
15. Depreciation, 10 % Fixed Capital/Yr
16. Taxes, 2% of Fixed Capital
17. Insurance, 1% of Fixed Capital
18. Other
19. TOTAL FIXED COST
TOTAL $
20. TOTAL OPERATING COST
21. COST: $/TonofCoal
22. Mill/kwh
6.34
3. 17
1,283,000
96,000
43,200
285,700
42,900
256,700
2,007,500
27,800
233,900
261,700
571,500
114,300
57, 100
742, 900
3,012, 100
42.60
9.49
1.42
8.52
66.65
0.92
7.?7
8.69
18.97
3.80
1.90
24.66
100.00
23. BY-PRODUCT CREDIT
(SEE FIGURE 4 )
-------
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1.38
1.15
1 .92
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\
2 3
*-$31/
v
\
/
ton
Maximu
EquivaU
of N in N
\ ,
\
m Price •
snt to vali
^-$60
s
\
je
/ton
^.
5 6
Basis: «120 megawatt power plant
•20 million cfh flue gas
•Sale of 83,000 tons of (NH4) S04 per year
Cost - Dollars Per Ton Coal
0.5 1.0 1.5 2.0 2.5
Cost - Mills per Kilowatt Hour
3.0
SHOWA-DENKO AMMONIACAL PROCESS : PROFITABILITY
-------
c. Commco Process
(1) Process Description^
The Cominco process involves the scrubbing of flue
gas with an aqueous solution of ammonium salts, and in this respect is similar to
the Fulham-Simon-Carves and Showa-Denko Ammoniacal processes. However,
in the Cominco process, the off-stream liquor is treated with sulfuric acid,
which results in the liberation of sulfur dioxide, and in the formation of ammonium
sulfate. The latter compound is isolated in the manner described in connection
with the other ammonia-based processes mentioned above. A flow diagram of the
Cominco process is shown in Figure 5.
(2) Process Reactions
Scrubber:
2 - » NHHSO
S02 -
(NH4)2S03 + 1/2 02 -
Acidifier:
2
(NH4)*2SO/+ H2S04" •> (NH4)2S04V SO-,f + H2O
(3) Chemical Requirements and By-Product Yields
Table 11 lists the raw material requirements and by-
product yields for treating one million SCF of flue gas by the Cominco process.
In addition, the annual requirements for three different sized plants, i. e. , 20
million SCFH, 0. 5 million SCFM, and 2. 5 million SCFM, respectively, are shown.
The raw material quantities allow for oxidation of 14% of the SO- to SO,, this
£» j
oxidation affects: (1) the quantity of ammonia needed to satisfy the reactions, and
(2) the relative quantities of the by-products, ammonium sulfate and SO_. The Hfo
621
value selected is the same as that used in the Zinc Oxide process evaluation and
is also essentially in agreement with the results of pilot-plant studies based on the
it1'i
Cominco process. Stack and in-plant losses have not been included in the data
shown in Table 11
-------
PURIFIED GAS
U)
NH,
FLUE
GAS
H20
SLUDGE
NH3
SCRUBBER (Port of/
Reactant Chemicals Per Million SCF Flue Gas Processed
A Stream
1
2
3
4
5
6
Component
so2
S°2
NH3
H2S04
( NH4) S04
so2z
Lb Moles
8.36
0.42
9.05
3.42
4.53
6.83
Lbs
535
26.8
154
335
597
437
LIOUEFIER
FILTER PRESS HEATER ACOtFIER EUMINATDR REACTION VACUUM CENTRIFUGE ORYER
TANK EVAPORATOR
COMINCO PROCESS : FLOW DIAGRAM
-------
TABLE 11
c
Raw Mate rials *a'
SO2 (flue gas)
NH3 (anhyd. )
H2S04 (100%)
by-Products
(NtLJoSO,
S02
COMINCO PROCESS:
HEMCAL REQUIREMENTS & BY-PRODUCT YIELDS
Quantity per Million Tons per Year
SC^ F^u.?,,5*.8, . ,, 20 Million 0 5 Million 2 5 Milli,
Lb Mole Lb SCFH Flue Gas SCFM Flue Gas SCFM Flue
7.94 508 40,230 60,350 301,730
9.05 154 12,195 18,290 91,460
3.41 335 26,530 39,800 198,980
4-52 597 47,280 70,920 354,600
6-83 437 34,610 51,960 259,800
(a)
-------
(4) Cost Estimate Detail
The cost estimates reported previously for the
Fulh,im-Simon-Carves and Showa-Denko Ammoniacal processes were based
on the assumption that the same basic equipment, with modifications as needed,
could be used for both processes. The Cominco process has been used primarily
on high SCX content (approximately 6%) waste gases from smelters and other
metallurgical operations. There are several variations of the process as used.
The basic process, however, consists of: (1) a dual or multi-scrubber system
for removal of SO2, (Z) conversion of ammonium bisulfite with sulfuric acid to
ammonium sulfate, sulfur dioxide, and water; and (3) stripping the residual SO,
with steam from the 40% ammonium sulfate solution. The stripped SO2 is used
to produce sulfuric acid in an adjacent plant, while the ammonium sulfate
solution is evaporated to produce crystalline material which can be used as a
fertilizer. A simplified flow diagram, Figure 6 , illustrates the process as
applied at The Consolidated Mining and Smelting Company of Canada, Limited,
Trail, B.C. An analysis of the process equipment indicated that the equipment
used m the Fulham-Simon-Carves process can be adapted, with minor changes,
to this process.
(a) Capital Costs
The equipment for the Cominco process to
handle removal of 95% of the SO- from flue gas in a 120, 000 kilowatt power plant
can be selected from the equipment used in the other ammonia-based processes.
For example, all of the Fulham-Simon-Carves process equipment can be used
with the exception of the autoclave. The following additional equipment would also
be needed:
• SO2 liquefaction system
• Acidifier
• Eliminator
An estimate of the purchased equipment cost follows:
-------
PURIFIED GAS
NH,
GAS
J L
;:t'V
X
_TE_
' f '*\
' 1 x
\
r
H20
i? HUH
i
T
SLUDGE T
SMELTER
STEAM
*
fc-Aj^ B=-
»^T
T
O
H2S04
{
-*-
STEAM^
1
1
7
AGIO PLANT
NH,
1
T
r
FERTILIZER PLAN
NH3
SCRUBBER
FILTER PRESS HEATER ACIOIFIER ELIMINATOR REACTION TANK
COMINCO PROCESS : FLOW DIAGRAM
(SMELTER GAS)
-------
Fulham-Simon-Carves installed equipment cost - $1,637,240
Autoclaves (2) -91,000 (installed)
SO2 liquefaction system +129. 500 (installed)
Revised installed equipment cost (1957): $1, 675, 740
Purchased equipment cost $1, 171,800
Current cost 1,317,000
Acidifiers (2) 30, 000 (purchased)
Eliminators (2) 50, OOP (purchased)
Current purchase cost: $1,397,000
The estimated total investment is $7, 278,400, or $60. 65/kw capacity. See
Table 12.
(b) Operating Costs
The estimated operating cost is summarized
in Table 13, which shows an annual cost of $3, 626, 900, or $7. 64/ton of coal and
3. 82 mill/kwh. Raw material costs were based on the quantities shown in Table 11.
Elimination of the autoclave and addition of
the other equipment was not expected to affect the labor requirement. Therefore,
both the direct labor and supervision are shown as equivalent to the Fulham-Simon-
Carves quantities, i.e. , 20 operators, 4 shift foremen, and one area superintendent.
Although the steam requirement was reduced
by 134, 700 M Ib per year due to elimination of the autoclave, additional steam
amounting to 22, 000 M Ib per year was needed in the eliminator to strip the
residual SO? in the liquor. The net reduction in steam compared with the Fulham-
Simon-Carves system was 112, 700 M Ib per year. It was also expected that the
SO7 liquefaction plant would add considerably to the utilities cost. A 3-4 stage
compression of SO?, requiring intercoolers and final condensers, was needed
194 647
to'produce anhydrous liquid SO-. Aries and Newton indicate a power
requirement of 0. 002 kwh and 9. 0 gallons of water per pound of SC^ produced.
The utilities cost shown in the Fulham-Simon-Carves process would be increased
by $830 per year in power costs and $31, 000 in water costs. Therefore, the total
utilities cost was $299, 700 for the Commco process.
-------
TABLE 12
COMINCO PROCESS: CAPITAL COST ESTIMATE SUMMARY
ITEM FACTOR COST - $
1. Purchased Equipment 1.00 1.397,000
2. Erection Labor O.Z5 349,300
3. Foundation b Platforms 0. 18 251,500
4. Piping 0.76 1.061.700
5. Instruments 0. 15 209.600
6. Insulation 0.08 111. 700
7. Electrical 0. 10 139. 700
8. Buildings 0.25 349,300
9. Land & Yard Improvements 0. 13 181, 600
10. Utilities 0.40 558, 700
11. PHYSICAL-PLANT COST 3.30 4,610, 100
12. Engineering It Construction p. 66 922, OOP
13. DIRECT PLANT COST 3.96 5,532, 100
14. Contractor's Fee 0. 19 265,400
15. Contingency 0.59 824, 300
16. FIXED CAPITAL COST 4. 74 6.621.800
17. Working Capital, 10% 0.47 656.600
TOTAL INVESTMENT 5.21 7.278.400
18. Capital Requirements
$/kw capacity 60. 65
-------
TABLE 13
COMINCO PROCESS: OPERATING COST ESTIMATE SUMMARY
Fixed Capital Cost: $6, 621. 800
ITEM
1. Raw Materials fc Chemicals
2. Direct Labor
3. Supervision
4. Maintenance, 5% of Fixed Capital
5. Supplies, 15% of Maintenance
6. Utilities
7. Other
8. TOTAL DIRECT COST
9. Payroll Burden, 20% of 2 It 3
10. Plant Overhead. 50% of 2, 3, 4 fc 5
11. Pack b Ship
12. Waste Disposal
13. Other
14. TOTAL INDIRECT COST
15. Depreciation, 10 % Fixed Capital/Yr
16. Taxes, 2% of Fixed Capital
17. Insurance, 1% of Fixed Capital
18. Other
19. TOTAL FIXED COST*
TOTAL $
1,617.800
120,000
43,200
331,100
49. 700
299. 700
2,461,500
32,600
272,000
304.600
662.200
132.400
66.200
860.800
44.60
3.31
1. 19
9. 13
1.37
8.27
67.87
0. 90
7.50
8.40
18.26
3.65
1.82
23.73
20. TOTAL. OPERATING COST
21. COST: $/Ton of Coal
22. Mill/kwh
7.64
3.82
23. BY-PRODUCT CREDIT
3.626.900
100.00
(SEE FIGURE?)
-------
The operating cost, without applying by-
product credits, was $7. 56 per ton of coal burned, which is somewhat higher
than the costs ol the other ammonia systems evaluated.
(c) Profitability
Figure 7 illustrates the levels of profit-
ability of this process with by-product sales at various unit prices, with an
expected cost of $4 -5 /ton of coal or 2-2. 5 mill/kwh.
d. Cominco Exorption Process
(1) Process Description
The Cominco Exorption process may be regarded ai
a variation of the Cominco process, in that the off-stream liquor is heated, rathe
than treated with sulfuric acid, for the liberation of sulfur dioxide. In principle,
no ammonium eulfate is produced, and no make-up ammonia is required.
However, some oxidation of sulfite to sulfate does occur, particularly during gas
scrubbing, and to this extent ammonium sulfate is also isolated. The process wi
used early in World War 11, when ammonia was in short supply. A flow diagram
of the Cominco Exorption process is shown in Figure 8.
(2) Process Reactions
Scrubber:
NH3 J
2TVTLT
wniij
1- H2O + SO2
k + H20 + S02 —
* NH4HSO3
—to (NH4)2S(
(NH4)2S03 + 1/2 02
Heater:
2NH4HS03
(3> Chemical Requirements and By-Product Yields
Table 14 shows the raw material needs and by-pr
production for the Cominco Exorption process.
-------
(NH4) S04 price - dollars per ton
CM
O
GO
O
Q>
O.
O
o
o>
u
31 25 20 IS 10 5 0
1 0
0.1 0
12345
Cost - Dollar's Per Ton Coal
—< i
Cost - Mills Per Kilowatt Hour
Basis: *120 megawatt power plant
•20 million cfh flue gas
•Sale of 47,280 tons (NH4> S04 and
34,610 tons S02 per year2
Current Prices:
• Liquid S02 : $69 / ton
• (NH) SO,: $31 /ton
COMINCO PROCESS : PROFITABILITY
-------
Readant Chemicals Per Million SCF Flue Gas Processed
PURIFIED GAS
A Stream
1
2
3
4
5
6
Component
S°2
S°2
NHo
1 NH4) $03
( NH/S04
so2 z
Lb Moles
8.36
a 42
2.22
6.83
Lll
6.83
Lbs
535
26.8
38.0
792
146
437
ro
SCRUBBER
SLUDGE
FILTER
PRESS
LIQUEFIER
FLASH TANK
VACUUM
EVAPORATOR
CENTRIFUGE
DRYER
COMINCO EXORPTION PROCESS : FLOW DIAGRAM
-------
TABLE 14
COMINCO EXORPTION PROCESS:
CHEMICAL REQUIREMENTS d BY-PRODUCT YIELDS
Quantity per Million Tons per Year
aw Materials**'
SO2 (flue gas)
NH3 (anhyd. )
iy- Products
'(NH4)2£04
so2
SCF Flue
Lb Mole
7.94
2.22
1. 11
6.83
Gas
Lb
508
38
147
437
70 Million
SCFH Flue Gas
40,230
3,010
11,640
34,610
0*5 Mil linn
SCFM Flue Gas
60, 350
4,520
17,460
51,960
2 5 Million
SCFM Flue Gas
301,730
22, 580
87. 300
259,800
^'Assumes 14% of SO2 oxidized to
-------
(4) Coat Estimate
In comparing the flow diagram, Figure 8 , with the
flow diagram of the Commco process, Figure 5 , it will be noted that minor
changes in equipment were made. It is also pointed out that the ammonium
sulfate system will be reduced in size since the quantity of ammonium sulfate
liquor handled and ammonium sulfate crystals recovered was substantially
reduced. The effect of this reduction is shown in the following paragraphs.
(a) Capital Costs
The purchased equipment cost of the Cominco
process was estimated at $1, 397,000. Several equipment changes were needed tc
convert the system to the Exorption process. The acidifier and eliminator
considered in the Cominco process would not be required in the Cominco Exorptio
process. However, additional required equipment would include two coolers, a
heater and flash tank, and a condenser. A reduction in the size of the vacuum
evaporator end centrifuge would also be applicable to the Exorption process. Sine
the daily usage of anhydrous ammonia is substantially reduced, the storage tank
was reduced from 229, 000 gallons to 35, 000 gallons. The net change amounted
to r$78, 300, resulting in a purchased equipment cost of $1, 475, 300. The
capital cost data are presented in Table 15, showing a total investment of
$7, 686, 300, or $64. 05/kw capacity.
(b) Operating Costs
Table 16 summarizes the operating costs for
the Exorption process. The annual cost amounts to $2, 726,400, or $5. 74/ton of
coal and 2. 87 mill/kwh. The raw material cost is much lower than in the Cominco
process since the absorbent is recirculated. Direct labor and supervision costs
are assumed to be the same as in the Cominco process.
The Cominco process data also provided the
basis for the utilities charges. The following changes were made:
-------
TABLE 15
COMINCO EXORPTION PROCESS: CAPITAL COST ESTIMATE SUMMARY
ITEM FACTOR COST - $
1. Purchased Equipment 1.00 1,475,300
2. Erection Labor 0.25 368,800
3. Foundation & Platforms 0. 18 265,600
4. Piping 0.76 1, 121,200
5. Instruments 0. 15 221,300
6. Insulation 0.08 118,000
7. Electrical 0. 10 147, 500
8. Buildings 0.25 368,800
9. Land & Yard Improvements 0. 13 191,800
10. Utilities p. 40 590.200
11. PHYSICAL-PLANT COST • 3.30 4.868. 500
12. Engineering b Construction 0. 66 973, 700
13. DIRECT PLANT COST 3.96 5.842,200
14. Contractor's Fee 0. 19 280, 300
15. Contingency 0.59 870,400
16. FIXED CAPITAL COST 4. 74 6.992, 900
17. Working Capital, 10% 0.47 693,400
TOTAL INVESTMENT 5.21 7,686, 300
18. Capital Requirements
$/kw capacity 64. 05
-------
TABLE 16
COMINCO EXORPTION PROCESS: OPERATING COST ESTIMATE SUMMAB
Fixed Capital Cost: $6, 992, 900
ITEM
TOTAL $
1.
2.
3.
4.
5.
6.
7.
8.
9.
10.
11.
12.
13.
14.
15.
16.
17.
18.
19.
20.
21.
22.
Raw Materials & Chemicals
Direct Labor
Supervision
Maintenance, 5% of Fixed Capital
Supplies, 15% of Maintenance
Utilities
Other
TOTAL DIRECT COST
Payroll Burden, 20% of 2 & 3
Plant Overhead, 50% of 2, 3, 4 & 5
Pick & Ship
Waste Disposal
Oiher
TOTAL INDIRECT COST
Dspreciation, 10 % Fixed Capital/Yr
Taxes, 2% of Fixed Capital
Insurance, 1% of Fixed Capital
Other
TOTAL FDCED COST
TOTAL OPERATING COST
COST: $/Ton of Coal 5. 74
Mill/kwh 2.87
180,600
120,000
43,200
349,600
52,400
756,400
-
1,502,200
32,600
282,600
_
-
-
315,200
699,300
139,800
69,900
_
909,000
2, 726,400
6.63
•™ ••»
4.40
•^ | nn
1.59
• ™ II • !»,
12.82
1.92
M^VMMM
27.74
••••••••WMV
55.10
IMMB0BM
1.20
••MMMMW
10.36
11.56
25.65
j.13
JZ.5J
•MMM
33.34
i •»• .1 •••"•
100.00
23. BY-PRODUCT CREDIT
(SEE FIGURE 9) _
-------
Steam:
Reduction of 17, 000 M Ibs due to the reduction in the size of vacuum
evaporator, ejectors and dryer.
Increase of 758,000 M Ibs for the heater requirement.
Net change: +741, 000 M Ibs.
Power:
Reduction of 477, 000 kwh due to reduction in size of centrifuge and
dryer.
Water:
i Reduction of 731, 000 M gal of circulating water due to reduction in
the capacity of barometric- and inter-condensers.
Increase of 2, 025, 000 M gal of circulating water for coolers and condenser.
Net change: +1, 294, 000 M gal.
The utility requirements thus became:
i
Steam: ' 1, 105, 000 M Ib per year
Power: 11, 759, 400 kwh per year
Makeup Water: 200, 000 M gal per year
Circulating Water: 2, 269, 000 M gal per year
(c) Profitability
Figure 9 illustrates the profitability of the
operation with the by-products sold at various prices. The anticipated cost is
approximately $4/ton of coal or 2 mill/kwh.
-------
( NH4) S04 price - dollars per ton
01234
Cost - Dollars Per Ton Coal
r
0
1 2
Cost-Mills Per Kilowatt Hour
Basis : • 120 megawatt power plant
•20 million cfh flue gas
• Sale of 11,640 tons (NH4) S04
and 34,610 tons S02 2
per year
Current Prices:
•Liquid S02 :$69/ton
•
-------
e. Zinc Oxide Process
(1) Process Description
In the Zinc Oxide process the flue gas is scrubbed
with an aqueous solution of sodium sulfite and sodium bisulfite. Zinc oxide is
mixed with the effluent liquor, forming insoluble zinc sulfite. This is filtered,
dried, and calcined to produce product sulfur dioxide and zinc oxide, which is
returned to the process.
Inasmuch as some oxidation occurs in the scrubber
to produce sulfate, which cannot be calcined, the process does include provisions
for its removal. The scrubber liquor is treated with insoluble calcium sulfite,
and the mixture is passed through a clarifier. The underflow from the clarifier,
which contains the calcium sulfite, is acidified with a portion of the product
sulfur dioxide, thereby causing the calcium sulfite to dissolve. Calcium ion is
thus made available for precipitation as calcium sulfate, which is removed by
filtration and discarded. The filtrate is treated with lime to precipitate calcium
sulfite, and it is then returned to the clarifier. A flow diagram for the Zinc
Oxide process is shown in Figure 10.
(2) Process Reactions
Scrubber:
2 NaHS03
Liming Tank:
2 NaHSO3 + CaO
Gasifier:
CaSO, + H-O + SO, »> Ca(HSO,),
' J C* £ 3 £*
Ca(HSO3)2 + Na2SO4 »- 2 NaHSO3 + CaSO4
Mixer:
2 NaHSO3 + ZnO
Calciner:
:S-
79
-------
Reactant Chemicals Per Million SCF Flue Gas Processed
PURIFIED GAS
A Stream
1
2
3
4
5
Component
S02
so2
CaO
CaS04-2H?0
S02 *
Ib-moles
8.36
a 42
Lll
Lll
6. S3
Ibs
535
26.8
62,2
191
437
CoO
FLUE GAS
SCRUBBER
LIMING TANK THICKENER MIXER DRYER
CLARIFIER GASIFIER FILTER
FILTER CALCINER
UQUEFIER
ZINC OXIDE PROCESS: FLOW DIAGRAM
-------
(3) Chemical Requirements and By-Product Yieldg
Table 17 indicates the raw material and by-product
requirements for treating one million SCF of flue gas by the Zinc Oxide process.
A complication which arises with processes in which SO_ is recovered as such is
£
that of partial oxidation of the SO? to sulfate in the gas scrubber. In the case of
594
the Zinc Oxide process this has been estimated by Johnstone and Singh as
occurring to the extent of about 10%, and by the Bureau of Mines report as
approximately 14%. The latter value was used for the present analysis.
Although the Zinc Oxide process is essentially
self-contained, various make-up chemicals are required as the result of spray
losses in the scrubber and dust losses in the calciner and drier. However,
this type of loss will, in general, occur in all of the processes under consider-
ation to about the same degree, and will consequently have little effect on the
relative economic ranking of the various processes. For this reason such
consistent losses were not considered in the present effort.
(4) Cost Estimate Detail
(a) Capital Costs
The capital cost estimates in the Bureau
of Mines report included the following:
Fixed Capital $2, 882, 550
Working Capital 288,260
Total Investment. $3,170,810
These values were adjusted as shown in the cost summary, Table 18, yielding
a total investment of $4, 430, 600, or $36. 92/kw capacity.
-------
TABLE 17
ZINC OXIDE PROCESS:
CHEMICAL REQUIREMENTS & BY-PRODUCT YIELDS
Quantity per Million
SCF Flue Gas
Lb Mole . Lb
Raw Materials
S02 (flue gas) 7. 94 508
CaO l.ll 62
By- Products
SO2 6.83 437
CaS04 1.11 191
Tons per Year
20 Million 0. 5 Million
SCFH Flue Gas SCFM Flue Gas
40,230 60,350
4, 890 7. 340
34,610 51,960
15,130 22,700
2. 5 Million
SCFM Flue G
301,730
36, 700
259,800
-------
TABLE 18
ZINC OXIDE PROCESS: CAPITAL COST ESTIMATE SUMMARY
ITEM FACTOR COST - S
1.
2.
3.
4.
5.
b.
^
6.
9.
iO.
11.
12.
13.
14.
15.
Ib.
IT.
if-
Purchased Equipment
Erection Labor
Foundation & Platforms
Piping
Instruments
insulation
Electrical
Buildings
Land & Yard Improvements
Utilities
Physical -Plant Cost
Engineering & Construction
Direct Plant Cost
Contractor's Fee
Contingency
Fixed Capital Cost
Working Capital. 10%
Tctal Investment
Capital Requirements
$/kw capacity 36.92
1. 00
0.25
0. 18
0. 76
0. 15
0.08
0. 10
0.25
0. 13
0.40
3. 30
0.66
3.96
0. 19
0.59
4. 74
Q.47
5.21
850,400
212.600
153,100
646,300
127,600
68,000
85,000
212,600
110,600
340/100
2.806,300
561,300
3,367,600
161,600
501,700
4,030.900
399. 700
4.430.600
-------
(b) Operating Costs
The estimated annual operating costs are
listed in Table 19, as $1, 874, 500, or $3. 95/ton of coal and 1. 98 mill/kwh.
Lime is the only raw material cost shown.
Direct labor amounts to 20 men. Super-
vision is the same as in previous processes, i. e. , four shift foremen and one
area superintendent. The utility requirements are as follows:
Steam: 228, 000 M Ib per year
Power: 3, 500, 000 kwh per year
Make-up Water: 634, 000 gallons
Fuel Oil: 3, 611, 000 gal per year
Approximately 22, 500 tons of wet calcium
sulfate cake contaminated with fly ash will be produced. Since it is doubtful
that this will have any value, a disposal operating cost should be considered.
On the basis of $4 per ton, a waste disposal cost of $90, 000 is included in the
operating cost.
(c) Profitability
Figure 11 shows the profitability of the
Zinc Oxide process as a function of SO, selling price. With liquid SO_ selling
£ £
for $20/ton the net cost would be about $2. 50/ton of coal or 1. 25 mill/kwh.
The cost of sodium carbonate and zinc oxide
losses are not reflected in the operating cost and profitability analysis. The
make-up quantities of these chemicals according to the Bureau of Mines report
could affect costs, as follows:
585,200 Ib/year ZnO = $89,200
2228 ton/year 58% Na_CO_ = 71, 300
& J • , .1
Added operating cost: $160, 500
-------
TABLE 19
ZINC OXIDE PROCESS: OPERATING COST ESTIMATE SUMMARY
Fixed Capital Cost: $4, 030, 900
ITEM
TOTAL $
1.
2.
3.
4
5.
*
7.
8.
9.
'10.
11.
12.
13.
14.
15.
16.
17.
18.
19.
20.
21.
22.
Raw Materials & Chemicals
Direct Labor
Supervision
Maintenance, 5% of Fixed Capital
Supplies, 15% of Maintenance
Utilities
Other
TOTAL DIRECT COST
Payroll Burden, 20% of 2 & 3
Plant Overhead, 50% of 2, 3, 4 &t 5
Pack & Ship
Waste Disposal
Other
TOTAL INDIRECT COST
Depreciation, 10 % Fixed Capital/Yr
Taxes, 2% of Fixed Capital
Insurance, 1% of Fixed Capital
Other
TOTAL FIXED COST
TOTAL OPERATING COST -
COST: $ /Ton of Coal 3.95
Mill/kwh 1.98
75,800
120,000
43,200
201,600
30,200
559,600
_
1,030,400
32,600
197,500
_
90,000
_
320, 100
403, 100
80,600
40,300
„
524,000
1.874, 500
4.04
6.40
2. 31
10. 7b
1.61
29.85
54.97
1.74
10.54
4.80
17.08
21. 50
4.30
2. 15
27.95
100.00
23. BY-PRODUCT CREDIT
(SEE FIGURE 11 )
-------
oo
™ 60
•o
J. 50
•i
fe 40
o.
12
CO
I 30
o>
^ 20
Q
\
\
Profit
i/urrer
\
\
Loss
IT f rice ^
\
\
K)V/ron
\
\
v
\
1012345
Cost - Dollars Per Ton Coal
i i i i i i i
0,5 0 0.5 LO 1.5 2.0 2.
Basis: -120 megawatt power plant
•20 million cfh flue gas
•Sale of 34,656 tons S02 per year
Cost - Mills Per Kilowatt Hour
ZINC OXIDE PROCESS : PROFITABILITY
-------
TABLE 20
HOWDEN-I.C.I. (CYCLIC LIME) PROCESS:
CHEMICAL REQUIREMENTS & BY-PRODUCT YIELDS
FOR THE SYSTEM USING LIME
Quantity per Million Tons per Year
SCF Flue Gas
w Materials
»
O_ (flue gas)
!aO
Lb Mole
7.94
7.94
Lb
508
445
20 Million
SCFH Flue Gas
40,230
35,215
0. 5 Million
SCFM Flue Gas
60, 350
52,825
2. 5 Million
SCFM Flue Gas
301, 730
264, 115
.-Products
;aSO3/CaSO4a 7.94 1302 103,120 154,680 773,390
Assumes 50% SO," oxidized to SO.".
alues shown based on the hydrated sulfite and sulfate, i. e. , CaSO,-2H?O and
CaSO4'2H2O
-------
TABLE 21
HOWDEN-I. C. I. (CYCLIC LIME) PROCESS:
CHEMICAL REQUIREMENTS & BY-PRODUCT YIELDS
FOR THE SYSTEM USING LIMESTONE
Quantity per Million Tons per Year
Raw Materials
SO, (flue gas)
b
CaCO3a
SCF Flue Gas
Lb Mole Lb
7. 94 508
10.72 1128
20 Million
SCFH Flue Gas
40,230
89,370
0. 5 Million
SCFM Flue Gas
60,350
134,055
2. 5 Million
SCFM Flue Ga
301,730
670, 340
By-Products
CaS03/CaS04b 7.94 1302 103,120 154,680 773,390
Assumes 35% excess of CaCO- of 95% purity.
Assumes 50% of SO ~ oxidized to SO.=
j 4
Values shown based on the hydrated sulfite and sulfate, i. e. , CaSO,- 2H O and
CaSO • ?H O 32
-------
As stated in reference 329, the amount of lime required
is essentially the stoichiometric amount, whereas a 35% excess of limestone
is used. In addition, the referenced report states that 30-80% of the calcium
sulfite is oxidized to the sulfate in the process. For purposes of this cost
estimate, a 50% oxidation was assumed.
(4) Cost Estimate
621
Bureau of Mines data were used as the basis
for this cost estimate.
(a) Capital Costs
The capital cost estimates in the Bureau
of Mines study are reproduced below:
Fixed c.ipital $1,811,100 %
Working capital 181, 100
Total Investment: $1, 922, 200
i
Up-dating these costs to November 1967 resulted in a total investment of
$2,659,700, or $22. 16/kw capacity, see Table 22.
(b) Operating Costs __-
Two operating cost estimates were prepared,
depending on whether lime or limestone was used as the raw material. These
estimates are presented as Tables 23 and 24, respectively. The costs are as
follows:
Lime Limestone
Total Operating Cost $1,996,500 $2,507,200
$/ton of coal 4.20 5.28
mill/kwh 2.10 2.64
The estimates are presented as Tables 23 and 24, respectively. The raw material
requirements are as follows:
Plant using lime: 35,215 tons of lime
*
Plant using limestone : 89, 370 tons of limestone
Direct labor was assumed to be the same
as in the Bureau of Mines study, i. e. , 4 shifts of 4 men per shift, or 16 men.
j.
A 35% excess of limestone was assumed.
-------
TABLE 22
HOWDEN-I.C.I. (CYCLIC LIME) PROCESS: CAPITAL COST ESTIMATE SUMMARY
ITEM FACTOR COST - $
1. Purchased Equipment 1.00 510'500
2. Erection Labor 0.25 127,600
3. Foundation & Platforms 0. 18 91.900
4. Pipmg 0.76 388,000
5. Instruments 0. 15 76,600
6. Insulation 0.08 40,800
7. Electrical 0. 10 51, 100
8. Buildings 0.25 127,600
9. Land & Yard Improvements 0. 13 66.400
10. Utilities 0.40 204,200
11. Physical-Plant Cost 3.30 1.684, 700
12. Engineering fe Construction 0. 66 336,900
13 Direct Plant Cost 3.96 2,021,600
14. Contractor's Fee 0.19 97,000
15. Contingency Q. 59 301,200
16. Fixed Capital Cost 4.74 2,419,800
17. Working Capital, 10% p. 47 239. 900
Total Investment 5.21 2.659.700
18. Capital Requirements
-------
TABLE 23
DEN-I.C.I. (CYCLIC LIME) PROCESS: OPERATING COST ESTIMATE SUMMARY
(Lime Used @ $15. 50/ton)
k
Fixed Capital Cost: $2, 419, 800
7.
8.
9.
0.
1.
2.
3.
4.
5.
6.
7.
8.
9.
1.
2.
ITEM
Raw Materials & Chemicals
Direct Labor
Supervision
Maintenance, 5% of Fixed Capital
Supplies, 15% of Maintenance
Utilities
Other
TOTA L DIRECT COST
Payroll Burden, 20% of 2 81 3
Plant Overhead, 50% of 2, 3, 4 & 5
Pack & Ship
Waste Disposal
Other
TOTAL INDIRECT COST
Depreciation, 10 % Fixed Capital/Yr
Taxes, 2% of Fixed Capital
Insurance, 1% of Fixed Capital
Other
TOTAL FIXED COST
TOTAL OPERATING COST
COST: $/Ton of Coal
Mill/kwh
4.20
TOTAL $
545.800
96,000
43,200
121,000
18.200
56.200
880.400
27.800
139.200
634.500
801.500
242.000
48.400
24.200
314.600
1.996.500
27.34
4.81
2. 16
6.06
0. 91
2.82
44. 10
1.39
6.97
31. 78
40. 16
12. 12
2.42
1,21
15.76
100.00
2. 10
3. BY-PRODUCT CREDIT
NOT APPLIED
-------
TABLE 24
HOWDEN-I.C.I. (CYCLIC LIME) PROCESS: OPERATING COST ESTIMATE SUMMARV
tLimestone Used @ $10/ton)
Fixed Capital Cost: $2,419, 800
ITEM
1. Raw Materials & Chemicals
2. Direct Labor
3. Supervision
4. Maintenance, 5% of Fixed Capital
5. Supplies, 15% of Maintenance
-------
Supervision requirements were assumed as one foreman per shift, and one area
supe rintendent.
The utility requirements were obtained u
factoring of the Bureau of Mines estimates, so as to apply to a 95% SO2 remot*'
rather than the 90% which was considered in the Bureau of Mines work.
Power: 7, 817, 000 kwh per year
Makeup Water: 93, 298, 000 gallons per year
For purposes of determining the tonnage
of waste materials generated by this system, it was assumed that 50% of the
calcium sulfite was oxidized to calcium sulfate. Accordingly, 103, 120 tons
of calcium sulfite/calcium sulfate are generated for the case where lime is
used. It was also assumed that the total weight of the waste would include
35% of free water. Therefore, the total weight of waste became 158, 640 tons.
In the case where limestone is used the
total weight of waste was 184, 000 tons, again on the basis that the cake contained
35% free water.
(c) Process Costs Using Limestone vs Lime
Although this estimate was made using
lime and limestone costs of $15. 50 and $10.00 per ton, respectively, it is
recognized that these figures may vary markedly depending upon availability
at any particular plant location. Thus the selection of either raw material
will ultimately depend on its delivered cost at the plant.
The use of cost values cited in a recent
180
article would have a very significant affect on process costs. In this paper,
in which a dry limestone process was compared with other dry processes,
the following unit costs were used:
Limestone: $2 per ton, delivered
Waste hauling and disposal: $0. 80 per ton, net
-------
These costs, applied to those given in Table 24, would reduce the raw matenalj
cost by $714, 700 and the waste disposal by $638, 100, resulting in a total
operating cost of $1. 154, 400. The unit cost in terms of dollars per ton of
coal would thus be reduced from $5. 28 to $2. 43. I here is a large variance
in reported limestone costs ($2 to $10 /ton) and in disposal costs. ($0. 80 to
$4. 00 /ton), which probably are Influenced by location and availability of the
limestone. The $10/ton cost of limestone used in this analysis was quoted
from representatives of the United Lime Division of Flintkote Company, Los
Angeles. The $4. 00 hauling and disposal cost used in the present study was
estimated from data given in Reference 57.
(d) Profitability
Since there are no by-productn of value
produced, credits cannot be applied to this process.
g. Batter sea Process
(1) Process Description
In the Batter sea process alkaline water is used foi
scrubbing sulfur dioxide from flue gas. The scrubbing capacity of such a medivt
is low, and large quantities of water are therefore required. The most suitable
water source is an alkaline river, which also serves as an acceptor for the
effluent water. Chalk or lime is added to the incoming water in order to in-
crease its alkalinity. The effluent is then treated with manganese sulfate and
air so that sulfate, rather than sulfite, can be returned to the river. The oxi-
dation is necessary in order to avoid any subsequent liberation of sulfur dioxide
downstream as a result of hydrolysis of the sulfite. A flow diagram of the
system is presented in Figure 13.
(2) Process Reactions
Scrubber:
or SO2 + CaO
Oxidizing Tank:
-------
PURIFIED GAS
vO
-J
FLUE GAS,
A Stream
1
2
3
4
5
6
7
Component
so2 _
so2
CaO
CaO
CaO
MnS04
CaS04
Lb Moles
8.36
0.42
6.35
1.59
7.94
0.018
7.94
Lbs
535
26.8
356
89
445
2.7
1080
or CoO.
IfTl
1
ALKALINE RIVER WATER
SPENT AIR
MnSO,
CoS04 SLURRY
TO RIVER
SCRUBBER
SLUDGE
FILTER SETTLER
MIXER
QXIDIZER
BATTERSEA PROCESS : FLOW DIAGRAM
-------
(3) Chemical Requirements and Bv-Product Yields
Table 2 5 lists the Taw material requirements and
by-product (waste) yields for treating one million SCF of flue gas by the
Battersea process. The SO^ calcium oxide and calcium sulfate quantities are
based on stoichiometric values. The water and manganese sulfate requirements
o C*7
are based on the data reported by Rees; the manganese sulfate requirement
was, in fact, increased in the present study in proportion to the greater SO2
content (0. 3%) of the standard flue gas as compared to that treated in the
Battersea plant. The waste mud quantities were also taken from this reference,
(4) Cost Estimate
The Battersea process is an effluent system in
that the scrubbing medium is used on a once-through basis and is discharged
into a waterway. Since the large quantities of water required limit its applica-
bility to select locations where water availability permits a noneffluent system,
the Howden-L C. L (Cyclic Lime) process was subsequently developed. (A detail-
ed discussion of the latter process is presented in the preceding section of this
report). Because of the fact that the two processes are similar and were both
developed by the English, the two have been compared to some extent on an
economic basis. Accordingly, cost-estimate data for the Howden-L C. I,
process was applied to the Battersea process when applicable.
(a) Capital Costs
108
Literature data indicate that the capital
investment for a plant which incorporates the Battersea process is approx-
imately two-thirds as great as the cost of a comparable sized plant which
employs the Howden - L C. L process. The purchased equipment cost of
the Howden - I. C. L process was estimated at $510, 500 (Table 22). The
purchased equipment cost for the Battersea process is therefore $340, 500.
Table 26 summarizes the capital costs. The total investment is $1, 774,100,
or $14. 78/kw capacity.
-------
TABLE 25
BATTERSEA PROCESS:
CHEMICAL REQUIREMENTS & BY-PRODUCT YIELDS
Quantity per Million
Tons per Year
SCF Flue Gas
Lb Mole Lb
Materials
i<. "" " '
>2 (flue gas) 7. 94 508
O Containing
200 ppm CaO
CaO (80% theor.) 6.35 356
H2O 10,890 196,000
.C03 (20% theor. ) 1.59 159
iS04 0. 06 9
Products
ad (wet solid waste) - 45
iSO4b 7. 94 1080
in effluent water)
20 Million
SCFH Flue Gas
40,230
28,200
15,523,200
(7830)a
12, 590
713
3, 565
85,540
OR Xyfi 11i f\Tt
SCFM Flue Gas
60,350
42,300
23,284,800
(11, 750)a
18,885
1,070
5,345
128, 310
2fi Vfillinn
SCFM Flue Gas
301,730
211,500
116,424,000
(58, 750)a
94,425
5,350
26, 730
641,550
Equivalent Water Rate, gpm
Values shown based on hydrated sulfate, i.e. , CaSO4'
-------
TABLE 26
BATTERSEA PROCESS: CAPITAL COST ESTIMATE SUMMARY
ITEM
1. Purchased Equipment
2. Erection Labor
3. Foundation & Platforms
4. Piping
5. Instruments
o. Insulation
7. Electrical
&. Buildings
9. Land fe Yard Improvements
iO. Utilities
11. Physical-Plant Cost
12. Engineering &t Construction
13 Direct Plant Cost
14. Contractor's Fee
15. Contingency
16. Fixed Capital Cost
17. Working Capital, 10%
Total Investment
18 Capital Requirements
$/kw capacity
FACTOR
COST - $
1.00
0.25
0. 18
0. 76
0.15
0.08
0. 10
0.25
0. 13
0.40
3. 30
0. 66
3.96
0. 19
Q. 59
4. 74
0.47
5.21
340,500
85, 100
61,300
258,800
51, 100
27,200
34, 100
85, 100
44, 300
136.200
1, 123,700
224, 700
1.348,400
64, 700
200.900
1.614.000
160, 100
1,774, 100
14. 78
-------
(b) Operating Costs
Two operating-cost summaries were
prepared. Table 27 illustrates a case in which alkaline water, similar to the
Thames Water, is used. The alkalinity of this water provides 80% of the
reactant for the SO in the flue gas. The remaining 20% reactant is provided
in the form of limestone. Table 28 lists the cost when chemically neutral water
is used and all of the limestone is added as slurry. The costs are as follows:
Thames Water Neutral Water
Total Operating Cost $851, 000 $1, 383, 000
$/ton of coal 1.80 2. 92
mill/kwh 0.90 1.46
Both direct labor and supervision were con-
sidered to be the same as in the Howden-I. C. I. process, i. e. , 4 operators and
1 foreman per shift and one area superintendent.
The utilities cost in this system is less
t lan in the Howden-I. C.I. process because it was assumed that water would
be available at no cost other than pumping charges which were included in
the power cost.
A relatively small charge for waste
disposal was included to handle the "mud" removed from the system.
Reference to operating cost, in terms
of dollars per ton of coal and mills per kwh, indicates a relatively inexpensive
system for SO_ removal'from flue gas, especially when using alkaline water.
(c) Profitability
Since there are no by-products of value
to recover, credits cannot be applied to this process.
-------
TABLE 27
BATTERSEA PROCESS: OPERATING COST ESTIMATE SUMMARY
(Thames River Water)
Fixed Capital Cost: $1, 614, 000
ITEM
1. Raw Materials & Chemicals
2. Direct Labor
3. Supervision
4. Maintenance, 5% of Fixed Capital
5. Supplies, 15% of Maintenance
-------
TABLE 28
BATTERSEA PROCESS: OPERATING COST ESTIMATE SUMMARY
(Neutral Water)
Fixed Capital Cost: $1, 614, 000
ITEM
1. Raw Materials b Chemicals
2. Direct Labor
3. Supervision
4. Maintenance, 5% of Fixed Capital
5. Supplies, 15% of Maintenance
l>. Utilities
7. Other
8. TOTAL DIRECT COST
9. Payroll Burden, 20% of 2 & 3
10. Plant Overhead, 50% of t, 3, 4 fc 5
11. Pack & Ship
12. Waste Disposal
13. Other
14. TOTAL INDIRECT COST
15, Depreciation, 10 % Fixed Capital/Yr
16. Taxes, 2% of Fixed Capital
17. Insurance, 1% of Fixed Capital
18. Other
19. TOTAL FIXED COST
20. TOTAL OPERATING COST
21. COST: $/Ton of Coal
22. Mill/lewh
2.92
TOTAL $
736,300
96,000
43,200
80,700
12.100
46,900
1,015,200
27.800
116,000
14,200
158,000
161,400
32,300
16,100
209,800
1, 383,000
53.24
6.94
3. 12
5.84
0.88
3.39
73.41
2.01
8.39
1.02
11.42
11.67
2.34
1.16
15. 17
100.00
1.46
>3. BY-PRODUCT CREDIT
NOT APPLIED
-------
h. Magnesium Hydroxide Process
(1) Process Description
In the Magnesium Hydroxide process, the SO2 is
scrubbed from the flue gas by reacting it with a slurry of magnesium hydroxide.41
The scrubber liquid which contains magnesium sulfite formed in the reaction is
aerated to oxidize the sulfite to sulfate. Synthetic ammonia is then added to
the circulating medium to regenerate the magnesium hydroxide.
A portion of the circulating medium is continuously
drawn off for product recovery. This side stream is filtered to remove the
solid magnesium hydroxide contained therein, vacuum evaporated, centrifuged,
and dried to recover the product ammonium sulfate. The magnesium hydroxide
precipitate from this side stream is returned to the circulating medium. A flow
diagram of the process is presented in Figure 14.
(2) Process Reactions
SO- Scrubber: >
Mg(OH)2 I + S02 » MgS03 i + H20
Oxidize r:
MgS03 i + 1/2 O2 •» MgSO4
Mixer:
MgSO4 + 2H20 + 2NH3 » Mg(OH)
(3) Chemical Requirements and By-Product Yields
Table 29 gives the chemical requirements and by-
product yields per milhon SCF of flue gas processed. Ammonia is the only
chemical requirement needed and ammonium sulfate is the only by-product.
(4) Fly-Ash Removal
The Magnesium Hydroxide process presented a special
problem as the result of the requirement for removing the fly ash entrained in the
flue gas. The problem was not a common one for the systems studied up to
-------
o
Ol
FLUE
GAST
PURIFIED GAS
A
',1\\
V!\\
Reactant Chemicals Per Million SCF Flue Gas Processed
A Stream
1
2
3
4
Component
so2
S°2
NH3
(NH4)S04
Lb Moles
8.36
0.42
15.9
7.94
Lte
535
218
270
1048
SPENT AIR
FLY ASH
FLY ASH FILTER
SCRUBBER
AIR-
S02 SCRUBBER
OXIOIZER
NMj MIXER FILTER
VACUUM CENTRIFUfl
EMKPORATOR
MAGNESIUM HYDROXIDE PROCESS: FLOW DIAGRAM
-------
TABLE 29
MAGNESIUM HYDROXIDE PROCESS:
CHEMICAL REQUIREMENTS & BY-PRODUCT YIELDS
Raw Materials
Quantity per Million
SCF Flue Gas
Tons per Year
20 Million
0. 5 Million
Lb Mole
2. 5 Millii
Lb .5CFH Flue Gas SCFM Flue Gas SCFM Flue
SO2 (flue gas) 7. 94
NH3 (anhyd.) 15.88
508
270
40,230
21,380
60,350
32,070
301,730
160,350
By-Product
(NH4)2S04
7-94
1048
83,000
124,500
622, 500
-------
this point because fly ash constituted the only particulate matter in the scrubbing
solutions and was easily removed by a simple filtration step. In this instance,
however, the scrubbing solution contains particulate magnesium sulfite and
magnesium hydroxide which cannot be discarded without introducing high
additional costs to the process.
One method which can be used to circumvent the
problem cited above is to acidify a portion of the circulating medium with sulfuric
acid to render the magnesium soluble, filter off the fly ash, and neutralize with
additional ammonia to recover the magnesium as magnesium hydroxide. Another
method involves incorporation of a fly ash scrubber similar in construction to the
SO, scrubber with the exception that only water would be circulated. The fly ash
would be removed by filtration of this circulating water.
Preliminary cost considerations indicated that re-
moval by filtration of the magnesium sulfite and magnesium hydroxide together
with the fly ash would increase the annual operating costs for the process by
about $11, 000, 000. The sulfuric acid approach would increase the cost by
approximately $6, 000, 000. On the other hand, it was estimated that annual
operating costs for the preliminary fly ash scrubber would be increased by about
$300,000, a relatively low cost. Clearly, the first two approaches mentioned
may be immediately rejected on the basis of the prohibitively high costs.
Accordingly, the fly ash scrubber was selected as the best method to achieve
fly ash removal, and it was incorporated in this cost estimate.
(5) Cost Estimate
Other than for the general description of the process
given in Reference 541, there was no data available regarding this process.
However, the system exhibited similarities to various phases of the Howden-L C. L ,
Fulham-Simon-Carves, and Zinc Oxide processes. Hence, the portions of these
three well-documented processes which are regarded as applicable to the re-
quirements of the Magnesium Hydroxide process were used as the basis for this
cost estimate.
-------
(a) Capital Costs
The fly-ash scrubber is the same as one
of the scrubbers used in the Fulham-Simon-Carves system. Choice of the
latter was based on the fact that the acidic conditions which would exist in
the fly ash scrubber as the result of SO2 solubility would present no problem
in the lead-lined system. A rotary vacuum filter required for the removal
of the fly ash from the system is incorporated in this subsystem. Cost: $273,40C
The Howden-I. C. I. system entails the use
of a lime or chalk slurry for removing SO2 from the flue gases. Because of
the similarity in the scrubbing medium in the Howden-I. C. I. process and
the Magnesium Hydroxide process, the scrubber assembly used for the former
was considered to be suitable for the latter process. Cost: $263, 600.
The air oxidation and mixer subsystems
were based on the use of two 150, 000 gal tanks of lead-lined steel construction.
An air compressor completes this system. Cost: $93,400.
The product recovery subsystem in the
Fulham-Simon-Carves process, with the exception of the autoclaves, was appliet
to this process. Cost: $428, 900.
Storage facilities were estimated to
cost $301,000.
Miscellaneous pumps resulted in an
additional cost of $32, 000.
The total purchased equipment cost
amounted to $1, 392, 300. The total capital investment of $7, 253, 900, or
$60. 53/kw capacity is summarized in Table 30.
(b) Operating Costs
The chemical requirements for the
Magnesium Hydroxide process are listed m Table 29. It is worthy of note
that anhydrous ammonia is the only raw material requirement and ammonium
sulfate is the only salable by-product.
-------
TABLE 30
MAGNESIUM HYDROXIDE PROCESS: CAPITAL COST ESTIMATE SUMMARY
ITEM FACTOR COST - $
1. Purchased Equipment 1.00 1,392. 300
2. Erection Labor 0.25 348,100
3. Foundation & Platforms 0.18 250,600
4. Piping 0.76 1,058. 100
5. Instruments 0.15 208,800
6. Insulation 0.08 111,400
7. Electrical 0.10 139,200
8. Buildings 0.25 348, 100
9. Land fc Yard Improvements 0. 13 181, OOP
10. Utilities 0.40 556,900
11. Physical-Plant Cost 3.30 4,594,600
12. Engineering fe Construction 0. 66 918, 900
13. Direct Plant Cost 3.96 5,513,500
14. Contractor's Fee 0. 19 264,500
15. Contingency 0.59 821. 500
16. Fixed Capital Cost 4. 74 6.599.500
17. Working Capital, 10% 0.47 654.400
Total Investment 5.21 7.253. 900
18. Capital Requirements
$/kw capacity 60. 45
-------
The estimated annual operating cost is
$3, 260, 100, or $6. 86/ton of coal and 3. 43 mill/kwh. See Table 31. The
cost for anhydrous ammonia was assumed to be $60 per ton.
Direct labor was assumed to be the same
as for the Fulham-Simon-Carves process in the Bureau of Mines study.
Supervision was assumed as one foreman per shift, and one area super-
intendent.
The utilities requirements for the
Magnesium Hydroxide process were based upon the data presented in the
Bureau of Mines report. A listing follows:
Steam: 229, 000 M Ib per year
Power: 10, 807, 000 kwh per year
Make-up Water: 200, 000 M gal per year
Circulating Water: 975, 000 M gal per year
(c) Profitability
Figure 15 indicates the profitability for the
Magnesium Hydroxide system, or a probable net cost of approximately
$4-4. 50/ton of coal and 2-2. 25 mill/kwh.
i. Magnesium Oxide Process
(1) Process Description
In the Magnesium Oxide process, the sulfur dioxide
is scrubbed from the flue gas by reacting it with a slurry containing excess
magnesium sulfite. The magnesium bisulfite obtained from this reaction is
then neutralized with magnesium oxide to regenerate the magnesium sulfite,
which is returned to the scrubber.
A portion of the regenerated magnesium sulfite
slurry is continuously drawn off for product recovery. The insoluble magnesium
sulfite is filtered, dried and calcined, resulting in by-product sulfur dioxide and
magnesium oxide, which is reused in the neutralization step.
As in the Zinc Oxide process, a complication arises
as the result of the partial oxidation of the sulfite to sulfate in both the calciner
-------
TABLE 31
MAGNESIUM HYDROXIDE PROCESS: OPERATING COST ESTIMATE SUMMARY
Fixed Capital Cost: $6, 599, 500
ITEM
1. Raw Materials & Chemicals
2. Direct Labor
3. Supervision
4. Maintenance, 5% of Fixed Capital
5. • Supplies, 15% of Maintenance
*>. Utilities
7. Other
8. TOTAL DIRECT COST
9. Payroll Burden, 20% of 2 & 3
10. Plant Overhead, 50% of 2, 3, 4 fc 5
11. Pack it Ship
12. Waste Disposal
13. Other
14. TOTAL INDIRECT COST
15. Depreciation, 10 % Fixed Capital/Yr
16. Taxes, 2% of Fixed Capital
17. Insurance, 1% of Fixed Capital
18. Other
19. TOTAL FIXED COST
TOTAL $
1.282.800
120.000
43,200
330,000
49.500
272,600
2,098, 100
32.600
271.400
304.OOP
660.000
132.000
66.000
858.000
39.35
3.68
1.33
10. 12
1.52
8.36
64.36
1.00
8.32
9.32
20.24
4.05
2.02
26.32
20. TOTAL OPERATING COST
3. 260. 100
100.00
21. COST: $/Ton of Coal
22. Mill/kwh
23. BY-PRODUCT CREDIT
6.86
3.43
(SEE FIGURE 15 )
-------
JL« \F9
1.61
= >•*
O>
I
§1.15
'E
£ a92
e
_w
f 0.69
£
k.
a.
0.46
0.23
0
*tu
35
30
cT
CO
CM
^ 25
I 20
a>
a.
S2
ID
1 15
1
s
i: 10
5
j
n
\
\
v
\
Curren
\
\
I price - '
\
\
SI /ton
V-Mw
>»
imum pr
ivalent to
in NH3~$
v
\
ice- I
value ofl
60 /ton
\
01234567
Basis : *120 megawatt power plant
•20 million rfh flue gas
Cost - Dollars Per Ton Coal
0.5 1.0 1.5 2.0 2.5
Cost - Mills Per Kilowatt Hour
3.0 3.5
Sale of 83, 000 tons (NH.) SO,
2
per year
MAGNESIUM HYDROXIDE PROCESS : PROFITABILITY
-------
and scrubber sections of the system. To prevent an e-fccessive buildup of sulfate,
the filtrate from the magnesium sulfite filtration is diverted to an evaporating
pond, from which magnesium sulfate is ultimately removed for disposal.
594
Johnstone and Singh indicate that approximately 10% of the absorbed sulfur
dioxide is oxidized in the scrubber. In addition, these workers estimated that
20% of the calcined magnesium sulfite would also undergo oxidation in the calciner.
Accordingly, it was assumed in the preparation of this cost estimate that 28% of
the absorbed sulfur dioxide will undergo oxidation. Recent developments, by
others, suggest that the degree of SO2 oxidation can be considerably less than
the value assumed herein; the effect of oxidation on process economics is quite
substantial.
A further process complication exists because of the
waste materials contained in the 95% pure magnesium oxide raw material.
Discussion with representatives of the Inorganic Chemicals Division of FMC
Corporation indicated that these waste materials consist of the oxides of cal-
cium, silicon, iron, and boron. These materials are expected to be insoluble
in basic media and would tend to build up in the system. Accordingly, these
wastes are removed from the system by discarding a portion of the precipitate
from the magnesium sulfite filtration. For the purpose of this cost estimate,
it was assumed that 5% of the precipitate is discarded. Fly ash is removed
in a prescrubber so that it does not recycle and build up in the magnesium
oxide slurry system. A flow diagram of the process is shown in Figure 16.
(2) Process Reactions
Scrubber:
MgS03 J + S02 + H20 » Mg(HS03)2
MgS03 f + 1/2 02 » MgS04
Neutralization Tank:
Mg(HS03)2 + MgO » 2 MgS03 f + H2 (80% of calcined MgSO-j)
MgSO3 + 1/2 O2 » MgSO4 (20% of calcined MgSC»3)
-------
Reactant Chemicals Per Million SCF Flue Gas Processed
A Stream
1
2
3
4
5
Component
SOg
so2
MgO
Mgso4
S02
Lb Moles
8.36
a 42
2.58
2.22
5.36
Lte
535
26.9
104
267
343
PURIFIED GAS
FLUE 6AS ,
WASTE
FLY ASH
FLY ASH FILTER
SCRUBBER PRESS
NEUTRALIZATION
SCRUBBER TANK
SLURRY
FILTER ORVER
CALCMER UQUEFIER
MAGNESIUM OXIDE PROCESS : FLOW DIAGRAM
-------
(3) Chemical Requirements and By-Product Yields
Table 32 indicates the chemical requirements and by-
product yields in treating one million SCF of flue gas by the Magnesium Oxide
process.
(4) Cost Estimate
The pulp and paper industry is currently using a
regenerative Magnesium Oxide process for pulp manufacture which is similar in
principle to the systems designed for removing SO, from flue gases. However,
598
the information reported in the literature is qualitative in nature and is of very
limited use as far as providing any basis for the present economic study is
concerned. One reference indicates that the process has been used in Russia for
the recovery of SO- from flue gases on a pilot-plant scale. However, no data
regarding the results were available.
Because of the similarities of the Magnesium Oxide
process with the processes investigated by Field, et al. , in the Bureau of Mines
study, the latter were used, wherever possible, as the basis of this estimate.
(a) Capital Costs
The same fly-ash scrubber subsystem
would be used for the Magnesium Oxide process that was used for the Magnesium
Hydroxide process. Cost: $273,400.
The same type of SO, scrubber in-
corporated in the Magnesium Hydroxide process was considered to be adequate
for the Magnesium Oxide system. Cost: $263, 600.
A 150, 000-gal tank was used in the
Howden-I. C.I. process to achieve a 3. 75-min. time delay. This tank size
and delay time is considered adequate to allow neutralization of the magnesium
bisulfite obtained as a product in the scrubber reaction. The material of con-
struction is lead-lined steel. Cost: $45, 700.
The waste filter having 200 square feet
of surface area was estimated to cost $16, 000.
-------
TABLE 32
MAGNESIUM OXIDE PROCESS:
CHEMICAL REQUIREMENTS AND BY-PRODUCT YIELDS
Quantity per Million Tons per Year
Raw Materials
SO2 (flue gas)
MgO
SCF Flue Gas
Lb Mole Lb
7. 94 508
2.58 104
_ o A Xyf-i 11i rtn
SCFH Flue Gas
40,230
8,670
0. 5 Million
SCFM Flue Gas
60, 350
13,005
2. 5 MUlii
SCFM Flue
301,730
65, OZ5
By-Product s
MgS04
SO,
2.22
5.36
267
343
21, 146
27, 165
31, 720
40, 748
158,600
203, 737
-------
A comparison of the material handling
requirements of the dryer and calciner sections of the Zinc Oxide system
cited in the Bureau of Mines report and the Magnesium Oxide process indicates
that, with the exception of the calciner operation, the requirements are
identical. Accordingly, the costs taken from the Zinc Oxide process study
for these items were used in the preparation of this estimate. The following
items of equipment were considered to be necessary: dryer, hammer mill,
cyclone separators, and blower at a total purchase cost of $79,200.
The calciner for the Magnesium Oxide
process would be used under markedly different operating conditions than that
used in the Zinc Oxide system. Effective calcination of zinc sulfite can be
achieved by heating the solid to approximately 850 F, whereas calcination
of magnesium sulfite requires that the solid be heated to approximately 1200 F.
It was considered that the calciner for the magnesium oxide would probably
be twice the cost of the unit used in the Zinc Oxide process. Cost: $248, 900.
The cost of the SO2 liquefaction system
was $107, 500. The cost of the slurry tank was estimated at $22, 200.
Various small pieces of equipment would
be required including pumps to handle the slurries and solutions, screw
conveyors to handle the solids, and the magnesium oxide and magnesium
sulfate storage systems. Based on the Bureau of Mines study, a cost of
$1J7, 500 was considered as adequate to cover these items.
i
From the estimated costs presented in
the foregoing, the total purchase price was determined to be $1, 174, 000.
The total investment was $6, 116, 600 or $50. 97 per kilowatt capacity. A
capital cost summary for the system is given in Table 33.
(b) Operating Costs
The annual requirement for magnesium
oxide is listed in Table 32 as 8, 670 tons. The OPD Reporter ( 5 February
1968) cites a price for calcined 95% magnesium oxide of $61. 00 per ton.
Accordingly, the annual cost of raw materials was determined to be $528, 900.
-------
TABLE 33
MAGNESIUM OXIDE PROCESS: CAPITAL COST ESTIMATE SUMMARY
ITEM
FACTOR COST - $
1. Purchased Equipment 1.00 1, 174.000
2. Erection Labor 0.25 293.500
3. Foundation fc Platform* 0. 18 211.300
4. Piping 0.76 892.200
5. Instruments 0. 15 176. 100
6. Insulation 0.08 93, 900
7. Electrical 0. 10 117.400
8. Buildingu 0.25 293,500
9. Land fc Yard Improvements 0. 13 152.600
10. Utilities 0.40 469.600
11. Physical- Plant Cost 3.30 3.874. 100
12. Engineexing fc Construction Q. 66 774.,800
13. Direct Plant Cost 3. 96 4. 648. 900
14. Contractor's Fee 0. 19 223. 100
15. Contingency 0. 59 692. 700
16. Fix«d Capital Cost 4. 74 5. 564. 700
17. Working Capital. 10% 0.47 551.800
Total Investment 5.21 6. 116.500
18. Capital Requirements
$/kw capacity _ 50.97
-------
The annual direct labor and supervision
was assumed to be the same as for the Magnesium Hydroxide process.
In the Bureau of Mines study, the utility
requirements are presented in itemized form for each of the three systems
considered. The present cost estimate for the Magnesium Oxide process
was largely based on equipment cited from this earlier study. A listing of
utility requirements follows. The requirement for fuel oil was obtained
from preliminary claculations of the heat requirements for calcining the
magnesium sulfite.
Power: 12, 107, 000 kwh per year
Steam: 228, 000 M Ib per year
Make-up Water: 200, 000 M gal per year
Circulating Water: 235, 000 M gal per year
Fuel Oil: 4, 472, 000 gal per year
The annual waste disposal cost was based
on the assumption that magnesium sulfate heptahydrate and 5% of the solids
collected in the magnesium sulfite filtration system would be discarded. Of
the total waste, impurities in the magnesium oxide amount to 434 tons. The
remaining waste, consisting of 6, 019 tons of magnesium sulfite hexahydrate
and 43, 306 tons of magnesium sulfate heptahydrate, would contain excess
water. Assuming 35% of the total weight of this waste is water, the total
overall waste amounts to 76, 550 tons. At $4. 00 per ton this became $306, 200.
A summary of the total operating costs
is presented in Table 34. This annual cost was estimated to be $2, 981, 500, or
$6.27/ton of coal and 3. 14 mill/kwh.
(c) Profitability
The OPD Reporter lists the current price
of magnesium sulfate as $2.45 per 100 Ibs. However, it is doubtful whether
large tonnages of this material could be sold at any price. From the meager
data available regarding price history, annual consumption, etc. , it appeared
unjustified at the present time to assume a credit for this material. An inde-
pendent study was made, however, in which the cost of isolation of magnesium
-------
TABLE 34
MAGNESIUM OXIDE PROCESS: OPERATING COST ESTIMATE SUMMARY
Fixed Capital Cost: $5, 564, 700
ITEM
1. Raw Materials & Chemicals
2. Direct Labor
3. Supervision
4. Maintenance, 5% of Fixed Capital
5. Supplies, 15% of Maintenance
6. Utilities
7. Other
8. TOTAL DIRECT COST
9. Payroll Burden, 20% of 2 & 3
10. Plant Overhead, 50% of 2, 3, 4 b 5
11. Pack b Ship
12. Waste Disposal
13. Other
14. TOTAL INDIRECT COST
15. Depreciation, 10 % Fixed Capital/Yr
16. Taxes, 2% of Fixed Capital
17. Insurance, 1% of Fixed Capital
18. Other
19. TOTAL FDCED COST
20. TOTAL OPERATING COST
21. COST: $/Ton of Coal
22. Mill/kwh
6.27
TOTAL $
528.900
120,000
43,200
278,200
41, 700
665,700
1,677,700
32,600
241,600
306,200
580,400
556,500
111.300
55,600
723,400
2,981.500
17.74
4.02
1.45
9.33
1.40
22.33
56.28
1.09
8.13
10.27
19.46
18.67
3.73
1.86
24.26
100.00
3.14
23. BY-PRODUCT CREDIT
(SEE FIGURE 17 )
-------
sulfate, via solvent removal, centnfugation, and drying, was determined as
an annual operating cost of $73, 700. On this basis the break-even point was
determined as $0. 17/100 Ibs; therefore, a process credit could be applied if
the material could be sold at a price higher than this.
The profitability of the Magnesium Oxide
process, based on the annual sale of 27, 165 tons of sulfur dioxide, and with no
credit applied for magnesium sulfate, is shown in Figure 17. This indicates
a. net cost of approximately $5 /ton of coal with liquid SO- being sold at $20 /ton.
j. Manganese Oxide Process
(1) Process Description
In the Manganese Oxide process, the sulfur dioxide
reacts with an aqueous suspension of manganese dioxide to form a solution of
manganese sulfate. Manganese dithionate and sulfuric acid are also formed
because of the presence of oxygen in the flue gas A portion of the absorbing
solution is continually withdrawn and allowed to settle The clarified portion
is then heated in an autoclave to precipitate the manganese sulfate which is
separated, dried, and calcined; the solid manganese oxide is then added to the
slurry from the clarifier and the mixture returned to the scrubber. The sulfur
dioxide which is released during the calcining step contains about 10% of sulfur
trioxide which is separated by liquefaction, converted to a sulfuric acid mist and
discharged as waste. A flow diagram of the process is shown in Figure 18.
(2) Process Reactions
Scrubber:
SO2 + MnO2 { - * MnSO4
3SO2 + 1/2 O2 + H2O + MnO2 f
Calcmer:
MnSO4 — £-»> Mixed Oxides (MnO2,
, MnO)
MnS,O, —-
cb A
-------
f VI
60
CO
S 50
3
zT
1 40
i_
S 30
•c 20
Q.
10
n
\
\
\
\
\
\
\
\
\
2345
Cost - Dollars Per Ton Coal
0.5 1.0 1.5 2.0 2.5
Cost-Mills Per Kilowatt Hour
Basis
3.0 3.5
• 120 megawatt power plant
•20 million cfh flue gas
*$3!e of 27,165 tons of SO^ psr ycsr
MAGNESIUM OXIDE PROCESS : PROFITABILITY
-------
tNi
LO
FLUE 6A8 ^
I to
FLT ASH
FLY ASH SCNUMER
FILTER
PUIIIFIED SAS
SO,
Readant Chemicals Per Million SCF Rue Gas Processed
A Stream
1
2
3
4
Component
*>z
s°2
*>2
S°3
LDMohs
8.36
a 42
7.15
a 79
Us
53$
26.9
457
63.2
UWEFICM
CLAftlFIEM AUTOCLAVE SCMMATM
CCNTRtFIMC
CALCIMER
MANGANESE OX IDE PROCESS : FLOW DIAGRAM
-------
Mixer:
MnS04 +
2 MhSO4
MhO + HS0 » MnS04 + HgO
(3) Chemical Requirements and By-Product Yields
Table 35 indicates the chemical requirements and
by-product yields on treating one million SCF of flue gas by the Manganese Oxide
process. The manganese oxides are completely recycled, and theoretically no
raw materials are needed.
(4) Cost Estimate
Development of the Manganese Oxide process
has been limited. Some pilot-plant studies have been conducted, however.
424
Tarbutton, et al. , conducted one investigation, and concluded from a
""""""" 439
similar method developed by the mining industry that the process is feasible.
Equipment similar to the processes investigated by Field, et al. , in the
Bureau of Mines study was used wherever possible as a basis for the cost
estimates in the present study.
(a) Capital Costs
In this process the fly ash in the flue
gas must be removed upstream of the SO- scrubber. Otherwise the ash
£*
would accumulate in the recirculated manganese oxide slurry from which
removal would be difficult. A single wood-grid packed, lead-lined steel
scrubber with a high water recirculating rate was selected for use as the
fly ash scrubber. The Fulham-Simon-Carves scrubber was considered
satisfactory for this purpose. Total purchase cost of unit: $231, 900.
The design of this tower was based
upon the data of Tarbutton which indicate a residence time of 7. 5 seconds
for complete removal of SO2 from flue gas, and Vedensky's439 experience
-------
TABLE 35
MANGANESE OXIDE PROCESS:
CHEMICAL REQUIREMENTS b BY-PRODUCT YIELDS
Quantity per Million Tons per Year
SCF Flue Gas 20 Million 0. 5 Million 2. 5 Million
Lb Mole Lb SCFH Flue Gas SCFM Flue Gas SCFM Flue Gas
,w Material
SO2 (flue gas) 7.94 508 40,230 60,350 301,730
'-Products
S02
S03 .
7.15
0.79
457
63.2
36, 190
4,910
54,280
7,370
271,420
36,850
-------
The cost of the 790 sq. ft. by 80 ft.
tall Fulham -Simon -Carve s scrubber was factored to obtain the cost of the
1, 119 sq. £t» by 80 ft. tall scrubber required here, resulting in a cost of
$358, 300.
A list of other equipment components
needed for this process follows:
Autoclave, stainless steel or Monel, 3800 gal, $30, 800
300 psi, 450°F
* V
CUrifiers (2), lead-lined steel 94, 000
Filters (4 including spares), vacuum rotary units 83,000
Centrifuge, vertical, perforated-basket, continuous 63,900
cone -type
Rotary dryer 46, 100
Calcine r, 1800°F operating temperature 230,000
Mix tank for MhOg slurry 8, 060
Swing hammer mill for dried MnSC*4 crystals 20, 200
Blower for calcine r 7, 230
Cyclone dust separator (2), one for SO, stream and 3, 320
one for flue gas from calciner unit
Screw conveyors, stainless steel 17, 140
Storage bin for MnSO, 29, 900
liquefaction unit 102, 000
$735,650
The total purchase price of all major
equipment for the Manganese Oxide process listed above is $1, 325, 850.
After factoring, the total investment cost becomes $6, 908,400 or $57. 56
per kilowatt capacity. A capital cost summary for the system is given in
-------
TABLE 36
MANGANESE OXIDE PROCESS: CAPITAL COST ESTIMATE SUMMARY
ITEM FACTOR COST * $
1. Purchased Equipment 1.00 1,326, OOP
2. Erection Labor 0.25 331,500
3. Foundation fc Platforms 0. 18 238. 700
4. Piping 0. 76 1.007,800
5. Instruments 0. 15 198,900
6. Insulation 0.08 106. 100
7. Electrical 0. 10 132.600
8. Buildings 0.25 331.500
9. Land fc Yard Improvements 0. 13 172.400
10. Utilities Q.40 530.400
11. Physical-Plant Cost 3.30 4.375.900
12. Engineering fe Construction Q. 66 875t 100
13. Direct Plant Cost 3.96 5.251.000
14. Contractor's Fee 0- 19 251,900
15. Contingency 0.59 782. 300
16. Fixed Capital Cost 4. 74 6,285,200
17. Working Capital, 10% 0.47 623,200
Total Investment 5.21 6.908.400
18. Capital Requirements
$/kw capacity 57.56
-------
(b) Operating Costs
The annual operating cost was estimated
at $2, 352, 500 equivalent to $4. 95/ton of coal and 2.48 mill/kwh. The list of
operating costs for the Manganese Oxide process is given in Table 37.
Theoretically, there is no make-up requirement for manganese oxide.
The annual direct labor and supervision
was a.ssumed to be the same as for the Zinc Oxide process.
The utilities for this process are:
Power: 8, 460, 000 kwh per year
Steam: 212, 000 M Ib per year
Make-up Water: 634, 000 M gal per year
Fuel Oil: 4, 960, 000 gal per year
(c) Profitability
The profitability of this system, based
on an annual sale of 35, 610 tons of liquid SO-, is given in Figure 19, which
indicates a net cost of $3. 50/ton of coal with liquid SO- selling for $20/ton.
k. Haenisch-Schroeder Process
(1) Process Description
183
In the Haenisch-Schroeder process, the sulfur
dioxide is scrubbed from the flue gas with water. Large quantities of water
are required because of the low solubility of the sulfur dioxide. Some lime
is continually added to the scrubber effluent in order to precipitate calcium sulfate;
formed through Oxidation of the sulfur dioxide. A flow diagram is shown in FiguN
-------
TABLE 37
MANGANESE OXIDE PROCESS: OPERATING COST ESTIMATE SUMMARY
Fixed Capital Cost: $6, 285, 200
ITEM
1. Raw Materials fc Chemicals
2. Direct Labor
3. Supervision
4. Maintenance, 5% of Fixed Capital
5. Supplies, 15% of Maintenance
f>. Utilities
7. Other
8. TOTAL DIRECT COST
9. Payroll Burden, 20% of 2 b 3
10. Plant Overhead, 50% of 2, 3, 4 fe 5
11. Pack & Ship
12. Waste Disposal
13. Other
14. TOTAL INDIRECT COST
15. Depreciation, 10 % Fixed Capital/Yr
16. Taxes, 2% of Fixed Capital
17. Insurance, 1% of Fixed Capital
18. Other
19. TOTAL FIXED COST
TOTAL $
120.000
43,200
314,200
47, 100
716,000
1,240,500
32,600
262,300
294,900
628.500
125.700
62.900
817.100
5.10
1.84
13.36
2.QO
30.43
52.73
1.39
11. 15
12.54
26.72
5.34
2.67
34. 73
20. TOTAL OPERATING COST
2.352.500
100.00
21. COST: $/Ton of Coal
22. Mill/kwh
23. BY-PRODUCT CREDIT
4. 95
2.48
(SEE FIGURE19)
-------
70
60
zT 50
40
-Current Price - $69 / ton
J2
£
I
3
•c
a.
30
20
10
\
Profit
\
Loss.
Basis: • 120 megawatt power plant
•20 million cfh flue gas
Sale of 35.610 tons S02 per year
012345
Cost - Dollars Per Ton Coal
0.5 0 0.5 1.0 1.5 2.0 2.5
Cost - Mills Per Kilowatt Hour
-------
AStream
1
2
3
4
Component
S02
502
S°2
CaO
Lb Moles
8.36
0.42
6.83
1.11
Lbs
535
26.8
437
62.2
PURIFIED GAS
FLUE GAS
SLUDGE
SCRUBBERS
CLARIFIER
FILTER
STRIPPER
REBOILER
TEAM LIQUID SO2
UQUEFIER
HAENISCH - SCHROEDER PROCESS
Figure 20
-------
(2) Process Reactions
Scrubber:
SO
H2O + SO2
Clarifier:
CaO + H2SO4 - •» CaSO4 t + H-,0
Stripper:
(3) Chemical Requirements and By-Product Yields
Table 38 gives the chemical requirements and
by-product yields in treating one million SCF of flue gas by the Haenisch-
Schroeder process. Because of the acidic nature of the scrubbing medium, it
is assumed that 14% of the SO2 is oxidized to sulfate. As was pointed out
in a previous section which covers the Zinc Oxide process, pilot-plant data have
shown that this degree of oxidation does indeed occur in an acid medium.
(4) Cost Estimate
The Haenisch-Schroeder process requires large
quantities of water, large -sized equipment, and a high utilities consumption.
(a) Capital Costs
The fixed capital cost for this process of
$29, 473, 000 is very high (see Table 39). Of the purchased equipment cost
($6,318,000), the heat-exchange equipment accounts for about two-thirds of
the total. A listing of all of the major equipment follows:
-------
TABLE 38
HAENISCH-SCHROEDER PROCESS:
CHEMICAL REQUIREMENTS & BY-PRODUCT YIELDS
Quantity per Million
SCF Flue Gas
Tons per Year
20 Million 0. 5 Million 2. 5 Million
Lb Mole Lb SCFH Flue Gas SCFM Flue Gas SCFM Flue Gas
Materials
(flue gas)
aO
7.94
1. 11
508
62.2
40,230
4,900
60,350
7,400
301,730
36,900
'roducts
aSOA
6.83
1. 11
437
151
34,610
11,960
51,960
17,900
259,800
89,500
-------
TABLE 39
HAENISCH-SCHROEDER PROCESS: CAPITAL COST ESTIMATE SUMMARY
ITEM
1. Purchased Equipment
2. Erection Labor
3. Foundation fc Platform*
4. Piping
5. Instruments
6. Insulation
7. Electrical
8. Buildings
9. Land & Yard Improvements
10. Utilities-
11. Physical-Plant Cost
12. Engineering fc Construction
13. Direct Plant Cost
14. Contractor's Fee
15. Contingency
16. ,F?xed Capital Cost
17. Working Capital, 10%
Total Investment
18. Capital Requirements
$/kw capacity
FACTOR
COST -$
1.00
0.25
0.18
0.76
0.15
0.08
0. 10
0.25
0.13
0.40
3.30
0.66
3.96
0.19
0.59
4.74
0.47
5.21
6, 218,000
I,554f500
1,119,200
4, 725, 800
932, 700
497, 400
621,800
1,554,500
808, 300
2,487,200
20, "519, 400
4. 103.900
24, (>23, 300
1, 181,400
3.b68.600_
29, 473, 000
2.922.000,
32.395.000,
269.96
-------
Reboiler for Stripper, 150, 000 sq. ft. surface
@ $4.71/sq. ft. $709,000
Heat Exchanger, SS tubes, 835, 000 sq. ft.
surface, @$3.50/sq. ft. 2,920,000
Cooler, S3 tubes, 124, 000 sq. ft. surface,
@ $4.65/sq. ft. 576,000
SO- Liquefier, 34,610 tons/year capacity 102,000
Scrubbing Tower, 85. 6 ft. x 85. 6 ft. x 50 ft. tall,
lead-lined steel 405, 000
Scrubbing Packing, 12 ft. depth of wood grids,
88,200 cu. ft. , @ $4. 31/ft3 381,000
Blowers (2, 1 spare), 20 MMSCFH, 2.5 in. H,O
static head, 125 hp . 32,700
Solution Pumps (8, 2 spares), 45,000 gpm each,
60ft. head, 900 hp, C.I. centrifugal, SS shaft 116,200
Fly Ash Clarifier, 135, 000 gpm overflow,
90,000 sq. ft. surface, 392ft. dia. ,
10 ft. deep, concrete 149, 000
Filter (2, 1 spare), continuous vacuum-type
rotary 41,500
' Stripping Column, 85. 6 ft. x 85. 6 ft. x 50 ft. tall,
lead-lined steel 405,000
Stripper Packing, 12 ft. depth of wood grids,
88,200cu. ft., @$4.31/ft3 381,000
Total purchase
equipment: $6,218,400
(b) Operating Costs
The estimated annual operating charges are
listed in Table 40.
The raw material cost is limited to the cost of
the lime used to precipitate the calcium sulfate.
Direct labor is assumed to be 4 men per shift
for 4 shifts. Supervision is assumed to be one foreman per shift and one
supervisor.
The utilities for the process are:
Steam: 11,820,000 M Ibs per year
Power: 33, 645, 000 kwh per year
Circulating Water: 35, 400, 000 M gal per year
-------
TABLE 40
HAENISCH-SCHROEDER PROCESS: OPERATING COST ESTIMATE SUMMARY
Fixed Capital Cost: $29, 473, 000
ITEM
1. Raw Materials fc Chemicals
2. Direct Labor
3. Supervision
4. Maintenance, 5% of Fixed Capital
5. Supplies, 15% of Maintenance
6. Utilities
7. Other
8. TOTAL DIRECT COST
9. Payroll Burden, 20% of 2 fe 3
10. Plant Overhead, 50% of 2, 3, 4 & 5
11. Pack & Ship
12. Waste Disposal
13. Other
14. TOTAL INDIRECT COST
15. Depreciation, 10 % Fixed Capital/Yr
16. Taxes, 2% of Fixed Capital
17. Insurance, 1% of Fixed Capital
18. Other
19. TOTAL FIXED COST
20. TOTAL OPERATING COST
21. COST: $/Ton of Coal 30.58
22. Mill/kwh 15.29
TOTAL $
76.000
72.OOP
43.200
1.473.500
221.000
7.882.000
9. 767. 700
23.000
904.800
927.800
2.947.000
589.400
294.700
3.831. 100
14.526.600
0.52
0.50
0.30
10.14
1.52
54.26
67.24
0.16
6.23
6.39
20.28
4.06
2.03
26.37
100.00
23. BY-PRODUCT CREDIT
(SEE FIGURE 21)
-------
The total operating cost for the Haenisch-Schroeder process is $14, 526,000 per
year or $30. 58 per ton of coal and 15. 29 mill/kwh.
(c) Profitability
The Haenisch-Schroeder process would operate
at a deficiency of $25. 56 per ton of coal burned if all of the SO- produced were
sold at $69/ton. The deficiency would be even greater (see Figure 21 ) if the SO,
had to be sold at lower prices.
1. Wet Thiogen Process
(1) Process Description
The Wet Thiogen process uses water as the
absorbent for sulfur dioxide. The scrubber effluent is treated with a solution
of barium sulfide to form sulfur and insoluble barium salts. The solids are
separated by settling and filtration, dried, and heated to distil off free sulfur.
The filtrate, essentially pure water, is returned to the scrubber. The still
resicue from the sulfur distillation, consisting primarily of barium sulfite and
sulfate, is then treated in a furnace to reduce the sulfur compounds to the
sulfide which is reused in the process. A flow diagram of the Wet Thiogen
process is shown in Figure 22.
(2) Process Reactions
Mixer:
2 BaS + 3 S0 - * BaSO + BaSO + 2 S
Heater:
BaS.,0, 45°-500 C . BaS03 + S
Furnace:
BaSO3 + 3 C - »• BaS + 3 CO
BaSO4 + 4 C - ••• BaS + 4 CO
-------
Current Price-$69/ton
TO
Basis
22 24 26 28
Cost - Dollars Per Ton Coal
10 11 12 13 14 15
Cost - Mills Per Kilowatt Hour
> 120 megawatt power plant
»20 million cfh flue gas
> Sale of 34,610 tons S02
per year
16
HAENISCH - SCHROEDER PROCESS : PROFITABILITY
-------
Reactant Chemicals Per Million SCF Flue Gas Processed
PURIFIED GAS
A Stream
1
2
3
4
5
Component
S02
SO,
C
CO
Lb Moles
8.36
a 42
7.94
17.0
17.0
Lbs
535
26.8
254
2(5
477
OJ
NO
SCRUBBER CLARIFIER FILTER MIXER CLARIFIER FILTER DRYER HEATER
WET THIOGEN PROCESS : FLOW DIAGRAM
-------
(3) Chemical Requirements and By-Product Yields
Table 41 gives the chemical requirements and the
by-product yield in treating one million SCF of flue gas by the Wet Thiogen pro-
cess. Coke is the only raw material consumed in the process and sulfur the only
by-product. A quantitative yield of sulfur is assumed since all oxidative products
are eventually reduced with coke. The barium sulfide is recycled. The carbon
monoxide is not considered as a salable by-product and is discharged through th6
furnace stack.
(4) Cost Estimate
The Wet Thiogen process was one of the earliest
methods for removing SO_ from smelter gases. Its development was limited,
£t
however, the process being abandoned after some pilot-plant work. The
absorption section, mixing tank, and solution clarifiers required for flue gas
applications are of very large size due to the low SO- solubility in water. The
remainder of the plant equipment dealing with sulfur recovery and recycle of the
barium sulfide was sized according to the quantity of sulfur handled.
For consistency with the many other processes
evaluated in this study, the cost of major equipment such as absorption columns,
pumps, filters, etc., were taken, whenever possible, from the Bureau of Mines
study reported in Reference 621. Other equipment costs were taken from
Reference 59.
(a) Capital Costs
The SO^ scrubber was sized to accommodate
the very high water-circulation rate required (135, 000 gpm). A lead-lined steel
tower with wood-grid packing similar to the Fulham-Simon-Carves tower is used.
With a liquid loading of 1100 gph/ft2, 7350 sq. ft. of tower cross section is
required. A 12-ft. depth of packing is used giving a 16-sec. gas-content time
in the scrubber.
-------
TABLE 41
WET THIOGEN PROCESS:
CHEMICAL REQUIREMENTS & BY-PRODUCT YIELDS
Quantity per Million Tons per Year
SCF Flue Gaa 20 Million 0. 5 Million 2. 5 Million
Lb Mole Lb SCFH Flue Gas SCFM Flue Gas SCFM Flue Gas
Materials
|iiD. (fiue gas) 7.94 508 40,230 60,350 301,730
b *
oke 17.04 204.5 16., 200 24,300 121,500
'roducts
7. 94
O 17.04
254. 1
477. 1
20, 120
37, 790
30, 180
56, 700
150,900
283,400
-------
The following scrubber section equipment
components are pertinent:
Tower, 85. 6 ft. by 85. 6 ft. , lead-lined steel, 50 ft. tall $405, 000
Packing, wood grids 88, ZOO ft3 at $4. 31 /ft3 381. 000
Blowers (2, 1 spare), 20 MM3CFH at 2.5 in. HgO static
head, 125 hp 32,700
Solution Pumps (4, 1 spare), 45, 000 gpm each at 60 ft.
head, 900 hp, CI centrifugal with SS shafts 58, 100
Fly Ash Clarifier, for 135, 000 gpm overflow, 90, 000
. ft. surface, rectangular or 392 ft. dia. , 10 ft.
, concrete
Fly Ash Filter (2, 1 spare), continuous, vacuum-type
rotary 41,500
Total purchase price: $1, 067, 300
Other major equipment is listed below:
Mixing Tank, 2-min. hold time, 214-ft. dia. by 10-ft. $74,500
deep, concrete
Solution Clarifier, 135, 000 gprn overflow, 392ft. dia.,
10 ft. deep, concrete 149, 000
Filters for barium salts and sulfur (3, 1 spare),
500 sq.
-------
After factoring, the total investment becomes $9, 065, 000 or $75. 54 per kilowatt
capacity. A capital cost summary for the system is given in Table 42.
(b) Operating Costs
The estimated annual operating costs f ^r th-:-
Wet Thiogen process are given in .Table 43. The annual cost for coke, the
major raw material, is $308,000.
Direct labor was assumed to be the same as
for the Fulham-Simon-Carves process, i. e. , 5 men per shift for 4 shifts.
Supervision was assumed as one foreman per shift and one area superintendent.
The utilities requirements for the Wet
Thiogen process are as follows:
Power: 18, 300, 000 kwh per year
Makeup Water: 93, 400 M gal per year
Circulating Water: 14, 250 M gal per year
Oil: 4, 280, 000 gal per year
The total operating cost amounts to $2, 916, 000 per year or $6. 14 per ton of coal and
3.07 mill/kwh.
(c) Profitability
This is a relatively high cost process. If
all of the sulfur produced is sold at $38/long ton, the net operating cost is still
an expensive $4. 71/ton of coal.
-------
TABLE 42
WET THIOGEN PROCESS: CAPITAL COST ESTIMATE SUMMARY
ITEM
1. Purchased Equipment
2. Erection Labor
3. Foundation fc Platform*
4. Piping
5. Instruments
6. Insulation
7. Electrical
8. Buildings
9. Land fc Yard Improvements
10. Utilities
11. Physical-Plant Cost
12. Engineering fc Construction
13. Direct Plant Cost
14. Contractor's Fee
15. Contingency
16. Fixed Capital Cost
17. Working Capital, 10%
Total Investment
18. Capital Requirements
$/kw capacity
FACTOR
COST -$
1.00
0.25
0.18
0.76
0.15
0.08
0.10
0.25
0.13
0.40
3.30
0.66
3.96
0.19
0.59
4.74
0.47
5.21
1,740,000
435,000
313,200
1,322,400
261.000
139,200
174,000
435,000
226, 200
696, 000
5.742,000
1^48.400
6,890,400
330, 600
1. 027. OOP
8,248,000
817.QOQ
9.0.6.5^00
75.54
-------
TABLE 43
WET THIOGEN PROCESS: OPERATING COST ESTIMATE SUMMARY
Fixed Capital Cost: $8, 248, 000
ITEM TOTAL $ %
1. Raw Materials fc Chemicals 308,000 10.56
25. ADJUSTED COST:
$/Ton of Coal 4. 71
Mill/kwh 2.36
145
2. Direct Labor 120,000 4.11
3. Supervision 43,200 1.48
4. Maintenance, 5% of Fixed Capital 412,000 14. 13
5. Supplies, 15% of Maintenance 61.900 2. 12
<>. Utilities 548, OOP 18.79
7. Other
8. TOTAL DIRECT COST 1,493,100 51.20
9. Payroll Burden, 20% of 2 & 3 32. 600 1. 12
10. Plant Overhead, 50% of 2, 3, 4 & 5 318. 500 IQ. 92
11. Pack b Ship . .
12. Waste Disposal - .,
13. Other
«§^SBMSJSBBBS^BBBBSBMS» MaMSMMBMSiaBBBB
14. TOTAL INDIRECT COST 351. 100 12.04
15. Depreciation, 10 % Fixed Capital/Yr 824.800 28.28
16. Taxes, 2% of Fixed Capital 164,900 5.65
17. Insurance, 1% of Fixed Capital 82.500 2. 83
18. Other
19. TOTAL FIXED COST 1.072,200 36.76
20. TOTAL OPERATING COST 2.916.400 100.00
21. COST: $/TonofCoal 6. 14
22. Mill/kwh 3.07
23. BY-PRODUCT CREDIT 680.400 -
-------
m. Ozone -Mn Ion and MnSO4 Processes
(1) Process Description
These processes involve the use of manganous ion
as an effective catalyst for the oxidation of sulfurous acid to sulfuric acid in
aqueous solution. With only the oxygen of the flue gas available as the oxidant,
the reaction is slow, and in the Ozone -Mn Ion process ozone is utilized to speed
up the reaction. Nevertheless, even under the best conditions, the reaction rate
is much slower than the neutralization of sulfur dioxide by the various basic
media used in other processes. With a synthetic flue gas, a retention time of
36 seconds for sulfur dioxide in the scrubber appears adequate with the Ozone -
Mn Ion process, but with a real flue gas the required time is more than doubled.
This may be due in part to the presence of trace quantities of phenols in the gas,
which tend to poison the catalyst. For either type of flue gas the retention times
would be much longer with the Manganese Sulfate process (without ozone). This
latter process will thus be much less attractive economically due to the larger
scrubber-design requirements. Therefore, it was not considered in the analysis
which is presented in the sections which follow,
Both processes are simple in operation, with the
product acid being directly obtained as a side stream from the scrubber.
However, the acid concentration never exceeds 40%, and consequently the product
has little economic value. A flow diagram of the Ozone -Mn Ion system is pre-
sented as Figure 23.
(2) Proces s Reactions
Scrubber:
(3) Chemical Requirements and By-Product Yields
Table 44 shows the raw material and by-product yield
requirements for treating one million SCF of flue gas by the Ozone -Mn Ion process.
The selected concentrations of chemicals, ozone at 160 ppm in the input gas and
Mn ion at 0. 3 g/100 g water in the scrubber liquor, are considered to be adequate
for the required degree of SO7 absorption; these values are based on the results of
424
pilot -plant work at TVA. Although sulfuric acid of approximately 30% strength
-------
PURIFIED GAS
HjO
Reactant Chemicals Per Million SCF Flue Gas Processed
A Stream
1
2
3
4
5
Component
s°2
S02
Ozone
MnS04
H2S04
Lb Moles
8.36
0.42
0.45
0.20
7.*
Lbs
535
26.8
21.6
29.7
778
FLUE GAS
AIR.
OZONE
XX XX
•MIST ELIMINATOR
MnS04 (CATALYST)
•»»ACID TO STORAGE
20 - 40%
OZONATOR SCRUBBER
SLUDGE
FILTER PRESS
OZONE -Mn ION PROCESS : FLOW DIAGRAM
-------
TABLE 44
Raw Materials
SO2 (flue gas)
MnSO,,
By-Products
(100%)
OZONE-Mn ION PROCESS
•
•
3HEMICAL REQUIREMENTS & BY-PRODUCT YIELDS
Quantity per
SCF Flue
Lb Mole
7.94
0.20
0.45
7.94
id) ICO ~
Million
Gas
Lb
508
29.7
21.6
778
1800
Tons per Year
20 Million 0. 5 Million
SCFH Flue Gas SCFM Flue Gas
40,230
2,350
1,700
61,620
142, 560
60.350
3,525
2,550
92,430
213,840
2.5 Millii
SCFMFlu;
301,731
17,62!
12, 750
462, 151
-------
is obtained as a direct by-product, the data in Table 44 lists the 100% acid.
Practical considerations indicate that concentration of the acid is perhaps not
feasible; nonetheless this optimistic route was taken in order to ascertain
whether the process, which is inherently quite unattractive, may have any
potential economic merit whatsoever. A subsequent discussion will elaborate
on this point.
(4) Cost Estimate
-^"""™-"— ^•••••••^^^•^
(a) Capital Costs
As mentioned previously, the high
retention or reaction time associated with the Ozone-Mi Ion process will result
in significant increases in the size and cost of the scrubber. In the Fulham-
Simon-Carves process, for example, the scrubbers were designed for a gas
velocity of 5 fps through the empty tower; the packing depth was only 6 ft. The
lead-lined steel construction of the scrubbers used in the Fulham-Simon-Carves
process is suitable to handle the 30-40% sulfuric acid produced in the Oaone-Mn
Ion process. Basic cost data therefore was obtained from the Bureau of Mines
study in which the Fulham-Simon-Carves process was considered.
In the referenced Bureau of Mines
report, the installed scrubber cost was equivalent to $2,, 66 per cubic foot.
i Correction to a purchased cost for 1967 gives a value of $2. 09/cu. ft. Due to
the huge size of the scrubbers required for a 36- to 88-sec. retention time, the
velocity through the empty scrubbers was reduced to 1 fps. This would require
a tower having a cross-sectional area of nearly 5, 600 ft. The total height of a
tower containing a 36-ft. column of packing was assumed to be 100 feet (gas
entry, liquid distributor, mist eliminator, and gas outlet). Cost of the tower,
based on the volume of the unit was calculated to be $1, 160, 000. The packing
and distributor costs were estimated in a similar manner, resulting in a value
of $837, 000. A tower of this size would probably have little utility. For example,
gas and liquid distribution would be a problem. The use of several towers used
in parallel and having smaller cross-sectional areas would increase the cost
additionally.
It was calculated that the pressure
drop in the large tower could be handled by a blower similar to those used in the
Fulham-Simon-Carves process. The present cost of two blowers (one stand-by)
was estimated at $33, 300.
-------
The cost of pumps for circulation of the
scrubbing liquor will be higher than might normally be expected due to the need
for Carpenter 20 or similar alloy material to insure compatibility with the 30%
sulfuric acid. An increase in pumping capacity is also needed because of the
high cross-sectional area of the towers. The estimated cost of the pumps was
$103,000.
The cost of two filter presses was estimated
at $57,000. A 150.000-gal lead-lined storage system for 30% acid was estimated
at $45, 700. The ozone was charged as a raw material in the ope rating-cost
estimate.
The total purchased equipment cost is
$2, 236, 000 for a system which would provide a 36-sec. retention time in the
absorbers. The total investment was estimated to be $11, 649, 500, or 97. 08/kw
capacity. See Table 45.
The purchased equipment cost for a system
providing an 88-sec. retention time will be substantially higher. If the scrubbing
tower and auxiliary equipment priced above are increased in size using the 0.6
factor, these costs become $3*410,000. The blower cost will increase by about
50% to $50, 000. All other equipment costs will be the same.
kii i
The total purchased equipment cost for a
system providing an 88-sec. retention is $3, 665, 700. It is evident that the
indicated fixed capital costs for both systems is excessive. However, the standard
^
factors used for auxiliary equipment may be high since the increase in scrubber
size will not necessarily increase the costs of all these components. The total
capital cost of $19, 098, 300 is summarized in Table 46.
I !
(b) Operating Costs
Operating costs for the 36-sec. System are
$3, 120, 500 per year or $6. 57/ton of coal and 3. 29 mill/kwh; for the 88-sec. system
the costs are $4. 586. 000 per year or $9. 65/ton of coal and 4. 83 mill/kwh. See
Tables 47 and 48.
The raw material quantities shown in Table 44
were used m the analysis. The cost of ozone in 1957 was established at $0. 105 per
424 i
lb, this value was used in the cost analysis although it is expected that the current,
-------
TABLE 45
OZONE-Mn ION PROCESS: CAPITAL COST ESTIMATE SUMMARY
(36-Sec. Reaction Time)
ITEM FACTOR COST - S
1. Purchased Equipment 1.00 2,236. OOP
2. Erection Labor 0.25 559,000
3. Foundation fc Platforms 0.18 402,500
4. Piping 0.76 1.699.400
5. Instruments 0.15 335,400
6. Insulation 0.08 178,900
7. Electrical 0.10 223,600
6. Buildings 0.25 559,000
9. Land & Yard Improvements 0. 13 290.600
10. Utilities 0.40 894,400
11. Physical-Plant Cost 3.30 7,378,800
12. Engineering & Construction 0. 66 1,475, 800
13. Direct Plant Cost 3.96 8,854,600
14. Contractor's Fee 0. 19 424,800
15. Contingency 0. 59 1,319,200
16. Fixed Capital Cost 4.74 10. 598,600
l~. Working Capital, 10% 0.47 1.050.900
Tctal Investment 5.21 11.649.500
if* Capital Requirements
$/kw capacity 97. 08
-------
TABLE 46
OZONE-Mn ION PROCESS: CAPITAL COST ESTIMATE SUMMARY
(88-Sec. Reaction Time)
ITEM FACTOR COST . $
1. Purchased Equipment 1.00 3.665.700
2. Erection Labor 0.25 916.400
3. Foundation & Platforms 0.18 659,800
4. Piping 0.76 2. 785,900
5. Instruments 0.15 549.900
6. Insulation 0.08 293,300
7. Electrical 0.10 366; 600
8. Buildings 0.25 916,400
9. Land & Yard Improvements 0.13 476,500
10. Utilities 0.40 1,466,300
11. Physical-Plant Cost 3.30 12,096,800
12. Engineering & Construction 0.66 2,419t400^
13. Direct Plant Cost 3.96 14.516,200
14. Contractor's Fee 0. 19 696.5QQ
15. Contingency Q. 59 2. 162.700
16. Fixed Capital Cost 4. 74 17.375.400^
17. Working Capital, 10% Q.47 1. 722.900^
Total Investment 5.21 19,098.300,
18. Capital Requirements
$/kw capacity 159. 15
-------
TABLE 47
OZONE-Mn ION PROCESS: OPERATING COST ESTIMATE SUMMARY
(36-Sec. Reaction Time)
Fixed Capital Cost: $10, 598, 600
ITEM
TOTAL $
1.
2.
3.
4.
5.
*.
7.
8.
9.
10.
11.
12.
13.
14.
15.
16.
17.
18.
19.
20.
21.
22.
Raw Materials & Chemicals
Direct Labor
Supervision
Maintenance, 5% of Fixed Capital
Supplies, 15% of Maintenance
Utilities
Other
TOTAL DIRECT COST
Payroll Burden, 20% of 2 & 3
Plant Overhead, 50% of 2, 3, 4 & 5
Pack & Ship
Waste Disposal
Other
TOTAL INDIRECT COST
Depreciation, 10 % Fixed Capital/Yr
Taxes, 2% of Fixed Capital
Insurance, 1% of Fixed Capital
Other
TOTAL FIXED COST
TOTAL OPERATING COST
COST: $/Ton of Coal 6.57
Mill/kwh 3.29
575,000
48,000
43,200
529, 900
79,500
98, 500
_
1,374, 100
18,200
350, 300
_
—
_
368,500
1,059,900
212.000
106.000
—
1.377.900
3. 120.500
'
18.43
1.54
1.38
16.98
2.54
3. 16
44.03
0.58
11.23
11.81
33.97
6.79
3.40
44.16
100.00
23. BY-PRODUCT CREDIT
NOT APPLIED
-------
TABLE 48
OZONE-Mn ION PROCESS: OPERATING COST ESTIMATE SUMMARY
(88-Sec. Reaction Time)
Fixed Capital Cost: $17, 375,400
ITEM
TOTAL $
1.
2.
3.
4.
5.
*,.
7.
8.
9.
10.
11.
12.
13.
14.
15.
16.
17.
18.
19.
20.
21.
22.
Raw Materials &c Chemicals
Direct Labor
Supervision
Maintenance, 5% of Fixed Capital
Supplies, 15% of Maintenance
Utilities
Other
TOTAL DIRECT COST
Payroll Burden, 20% of 2 & 3
Plant Overhead, 50% of 2, 3, 4 & 5
Pack & Ship
Waste Disposal
Other
TOTAL INDIRECT COST
Depreciation, 10 % Fixed Capital/Yr
Taxes, ?,% of Fixed Capital
Insurance, 1% of Fixed Capital
Other
TOTAL FIXED COST
TOTAL OPERATING COST
COST: $/Ton of Coal g. ^5
Mill/kwh 4.83
575,000
48,000
43,200
868,800
1 130,300
98, 500
.
1,763,800
18,200
545, 200
.
-
.
563,400
1,737,500
347, 500
173,800
.
2.258,800
4,586,000
12.54
1.05
0.94
18.94
2.84
2.15
4,
38.46
0.40
11.88
-
-
•
12.28
37.89
7.58
3.79
•>
49.26
100.00
23. BY-PRODUCT CREDIT
(SEE FIGURE 24)
-------
Direct labor was estimated at a relatively
low level of two men per shift. Supervision of one foreman per shift plus one
area supervisor was considered as adequate.
The utility requirements for both variations
of this process are as follows:
Power: 1, 550, 000 kwh per year
Makeup Water: 34, 200 M gal per year
Circulating Water: 47, 500 M gal per year
The operating cost for the system with the
36-sec. reaction time is presented only as a hypothetical case which is not
achievable with real flue gas. The operating cost of $9. 65 per ton of coal
(without by-product credit) for the system having the 88-sec. reaction time
was used for comparison with the economic considerations developed for the
other processes.
(c) Profitability
The data shown in Table 44 indicate that a
substantial quantity of sulfuric acid (approximately 30%) is recoverable as a
by-product. It must, however, be concentrated to at least 78% to be commercially
valuable. If it were possible to recover the value of this acid (as 100%) at
current prices, it would yield an income of $2, 058, 100. This would reduce
operating costs by $4. 33 per ton of coal burned. Subtracting this from the $9. 65
operating cost shown in Table 48 yields a net operating cost of $5. 32 per ton of
coal; this value is still very high when compared with the results for some of the
other processes.
These cost values would be even higher if the
cost of concentrating the acid were included. In a comprehensive study of the
method, Johnstone found that acid of strength not exceeding 4% is produced
when flue gas from coal combustion is used. Accordingly, the results pre-
sented in the foregoing are very optimistic. Because the process is economically
uncompetitive, as considered above, no further attempt was made to quantitatively
determine what the effect of producing the 4% acid rather than the 30 or 40%
product would have on process costs; obviously, the costs would be even higher.
Figure 24 shows the optimistic profitability of the process for the conditions
specified.
-------
tn
50
~ 40
CM
30
g 20
i
3 10
Current Price : $33.40 per ton
l
45678
Cost - Dollars Per Ton Coal
i 1 i
234
Cost - Mills Per Kilowatt Hour
Basis: *120 megawatt power plant
•20 mi I lion cfh fiuegas
•Sale of 61,620 tons
(100%) per year
10
OZONE - Mn ION PROCESS : PROFITABILITY
-------
n. Sulfidine Process
(1) Process Description
In the Sulfidine process a mixture of xylidine and
ter (approximately 1:1) is used as the absorbent for sulfur dioxide. The
sorbent is then stripped of the gas by the use of indirect steam and pure
fur dioxide is recovered after washing, drying, and compression. The
>cess is complicated by the necessity of scrubbing xylidine vapors from the
dte gas by the use of dilute sulfuric acid in a separate scrubbing tower. The
idine sulfite solution must then be stripped to remove the xylidine. A further
nplication of the process is the necessity of adding soda ash to the absorption
stem in order to control the formation of sulfates formed from the oxidation of
.fur dioxide. The resulting sodium sulfate is removed from the system with
• waste water stream. A flow diagram of the sulfidine process is given in
gure 25.
(2) Process Reactions
Scrubber:
Stripper:
(XH) HSO3 - to X + SO2 + H2O
(XH)2SO3 - to 2 X + SO2 + H2O
(3) Chemical Requirements and By-Product Yields
Table 49 gives the raw material requirements and
-product yields for treating one million SCF of flue gas by the Sulfidine process.
234
this process, the extent of SO- oxidation in solution was reported by Katz to
one percent for the case of a smelter gas containing 5% SO2. However,
= mixture of xylidine isomers, C^H
-------
Rsadant Chemicals Per Million SCF Flue Gas Processed
PURIFIED GAS
DILUTE
A Stream
1
2
3
4
5
Component
S°2
S°2
S02
HoSOj
Soda Ash
Lb Moles
8.36
a 42
6,83
L39
2.22
Lbs
535
26.8
437
136
235
OB
XYLIOINE
STMPPER
RE-BOILER
XYLIOINE RECOVERY
SCRUBBER FILTER PRESS SCRUBBER UNIT SEPARATOR
SULFIDINE PROCESS : FLOW DIAGRAM
Figure 25
-------
TABLE 49
SULFIDINE PROCESS:
CHEMICAL REQUIREMENTS & BY-PRODUCT YIELDS
Quantity per Million Tons per Year
SCF Flue Gas
.Materials
i
P2 (flue gas)
!2SO4 (100%)
1
yhdine,
Lb Mole
7.94
1.39
2.22
0. 165
Lb
508
136
235
20
20 Million
SCFH Flue Gas
40,230
10, 740
18,610
1,580
0. 5 Million
SCFM Flue Gas
60,350
16, 110
27,915
2,370
2. 5 Million
SCFM Flue Gas
301,730
80,550
139,580
11,350
jj2NH2
3roducts
a2S04
6.83
2.22
437
595
34,610
47, 120
51,960
70,680
259,800
353,400
-------
Flemmg145found that the oxidation was much higher for low concentration SO2
gases. Literature data672indicate that the extent of oxidation is essentially
independent of the initial SO2 concentration in the gas. In order to be consistent
with the Zinc Oxide process, which was discussed in a previous section, and
for which oxidation data are available, 14% of the SO2 absorbed was assumed to be
oxidized in this process. The loss of xylidine as vapor from the scrubbing column
and in the various liquid streams had to be considered in this process, since it is
a significant item of cost. Data for plants treating 4 or 5% SO2 gas were extrapo-
lated to a flue gas containing 0. 3% SO2- The xylidine loss of 20 Ib/MMSCF flue gas
although quite high, is believed to be conservative for this process. The chemical
requirements are based on an absorption temperature of 30 C.
(4) Cost Estimate
In order to be as consistent as possible with the cost
estimates conducted for the other processes, a continuing effort was made to
utilize the same cost bases to the fullest extent possible. In this particular case,
most of the equipment costs for the Sulfidme process were based on the Bureau of
Mines cost estimates for the Fulham-Simon-Carves, the Zinc Oxide, and the
Howden-I. C.I. (Cyclic Lime) processes. Equipment costs shown herein were
corrected to 1%7 purchase costs and adjusted for size according to the 0.6 power
factor.
(a) Capital Costs
The fixed capital cost for this process of
$10, 512, 900 is high (see Jable 50 ). Of the total purchased equipment cost of
$2, 217, 900, more than half ($1, 227, 000) reflects the cost of the three main gas
scrubbers, packing, and liquid pumps. These large columns, with a cross-
sectional area greater than 2, 200 square feet, have a liquid capacity of 36, 000 gpm.
A moderately slow gas-absorption rate (about 10 seconds contact time) further
contributes to chese large dimensions. Lead-lined steel towers with wood packings
similar to the tower used in the Fulham-Simon-Carves process were considered.
The heat exchanger cost of $478, 000 constitutes another major cost item; this
exchanger of Type 316 stainless steel is required to transfer heat to the 1, 750 gpm
of corrosive SO2-nch solution from the hot, lean, stripping column solution. The
cost of the SO2 liquefaction system is $102, 000.
-------
TABLE 50
SULFIDINE PROCESS: CAPITAL COST ESTIMATE SUMMARY
ITEM FACTOR COST - $
1. Purchased Equipment 1.00 2, 217, 900
2. Erection Labor 0.25 554. 500
3. Foundation fc Platform* 0. 18 399.200
4. Piping 0.76 1.685.600
5. Instruments 0. 15 332. 700
6. Insulation 0.08 177,400
7. Electrical 0. 10 221.800
8. Buildings 0.25 554. 500
9. Land & Yard Improvements 0. 13 288, 300
10. Utilities Q.40 887.200
11. Physical- Plant Cost 3.30 7.319. 100
12. Engineering b Construction Q. 66
13. Direct Plant Cost 3.96 8. 782. 900
14. Contractor's Fee 0. 19 421.400
15. Contingency 0.59 1. 308.600
16. Fixed Capital Cost 4. 74 10.512.900
17. Working Capital, 10% 0.47 1.042.400
T otal Inve s tment 5.21 1 1.555. 300
18. Capital Requirements
$/kw capacity 96.29
-------
(b) Operating Costs
The estimated annual operating charges are
listed in Table 51. The total operating cost of $5, 803, 000 per year, or
$12.22 per ton of coal, is quite high when compared with operating costs of
ether processes. The largest single item contributing to cost is xylidine,
commercially available in tonnage lots as a mixture of the isomers.
The cost of utilities is also relatively high
with steam being the most expensive item. This high cost is due to the large
molar ratio of steam in SO2 (30/1) used in the stripping column. The fact that
the SO2 content in the inlet flue gas (0. 3%) is low. combined with the need of
using large quantities of scrubbing solution, results in a low SO, concentration
(12. 9 g/liter) in the rich solution. These considerations, in addition to the fact
that the SO- vapor pressure of the xylidine-water solution does not change
£ _
significantly as the temperature is raised from 30 to 100 C, accounts for the
large steam requirement. The utilities for the SO- liquefaction system was
estimated at $31, 800. Disposal cost of the waste sodium sulfate was based on
the weight of the crystalline material.
(c) Profitability
The Sulfidine process would operate at a deficit
of $7. 19 per ten of coal burned (see Figure 26 ), if all of the SO, produced were
L»
scld for $69/ton. If the more realistic price of $19/ton (equivalent sulfur value)
was used for all the SO2 produced, the process would operate at a cost of
$10. 75/ton of coal burned. In calculating the above costs, no charges were made
to cool the inlet gas to 30 C. In all other processes the gas absorption was
assumed to take place at 50°C. This additional cooling would take additional
heat-transfer surface and cooling water and further increase the plant costs.
-------
TABLE 51
SULFIDINE PROCESS: OPERATING COST ESTIMATE SUMMARY
Fixed Capital Cost: $10, 512, 900
ITEM
TOTAL $
1.
2.
3.
4.
5.
7.
8.
9.
10.
11.
12.
13.
14.
15.
16.
17.
18.
19.
20.
21.
22.
Raw Materials fe Chemicals
Direct Labor
Supervision
Maintenance, 5% of Fixed Capital
Supplies, 15% of Maintenance
Utilities
Other
TOTAL DIRECT COST
Payroll Burden, 20% of 2 & 3
Plant Overhead, 50% of 2, 3, 4 fe 5
Pack & Ship
Waste Disposal
Other
TOTAL INDIRECT COST
Depreciation, 10 % Fixed Capital /Yr
Taxes, 2% of Fixed Capital
Insurance, 1% of Fixed Capital
Other
TOTAL FIXED COST
TOTAL OPERATING COST
COST: $/ Ton of Coal 12.22
Mill/kwh 6.11
2,422,000
120,000
43,200
525.600
78.800
641.800
_
3.831,400
3?.. 600
383.800
—
188.500
—
604, 900
1,051,300
210, 300
105, 100
_
1,366,700
5,803,000
41.73
2.07
0.74
9.06
1.36
11.06
_
66.02
0.56
6.62
—
3.25
—
10.43
18. 12
3.62
1.81
_
23.55
100.00
23. BY-PRODUCT CREDIT
(SEE FIGURE 26)
-------
cr-
"»
o»
a.
to
I
g
70
60
50
40
30
20
10
Current Price: $69
3.5 4.0 4.5 5.0 5.5 6.0
Cost - Mills Per Kilowatt Hour
Basis: • 120 megawatt power plant
•20 million cfh flue gas
•Sale of 34,610 tons S02 per year
8 9 10 11 12 13
Cost ~ Dollars Per Ton Coal
6.5
SULFID1NE PROCESS : PROFITABILITY
-------
o. Basic Aluminum Sulfate Process
(1) Process Description
The Basic Aluminum Sulfate process involves the
absorption of sulfur dioxide with a solution of a soluble basic aluminum salt,
followed by thermal stripping of the sulfur dioxide and recycling of the re-
generated scrubber liquor. The choice of the absorbing medium is based on
the ease with which it can be regenerated in the stripping step. In common
w.th other regenerative methods, oxidation of sulfite to sulfate represents a
process complication. In the present process, sulfate is removed as insoluble
calcium salfate through treatment with limestone. A flow diagram of the Basic
Aluminum Sulfate process is shown in Figure 27.
(2) Process Reactions
Scrubber:
A1(OH)SO4 + SO2 »• A1(OSO2H)SO4
A1(OSO2H)S04 + 1/2 O2 » A1(OSO3H)SO4
Stripper:
A1(OS02H)SO4 » A1(OH)SO4 + SO2 {
Mixing Vessel:
CaCO3 + A1{OSO3H)SO4 »> CaSO4 ( + CO2
A1(OH)S04
(3) Chemical Requirements and By-Product Yields
Table 52 gives the raw material and by-product data
:-r treat-rig one million SCF of flue gas by the Basic Aluminum Sulfate process.
Ir. smelter-plant applications of this process, 1 to 1-1/2% of the SO2 absorbed
irom a 5% feed gas is oxidized to sulfate in solution. As discussed previously
i- cor junction with the Sulfidine process, however, the extent of oxidation is
much higher for low concentration SCX gases. In order to be consistent with
several other processes for which oxidation data are available, 14% of the SO-
absorbed is assumed to be oxidized to sulfate.
-------
Reactant Chemicals Per Million SCF Flue Gas Processed
A Stream
1
2
3
4
5
Component
so2
SG2
CaS04-2H20
Lb Moles
8.36
0.42
6.83
Ul
1.11
Lbs
535
26.8
437
115
190
PURIFIED GAS
cr-
SLUDGE
SCRUBBER FILTER PRESS
COOLER
FILTER PRESS
MIXIN6
VESSEL MIXER
BASIC ALUMINUM SULFATE PROCESS : FLOW DIAGRAM
-------
TABLE 52
BASIC ALUMINUM SULFATE PROCESS:
CHEMICAL REQUIREMENTS AND BY-PRODUCT YIELDS
Quantity per Million Tons p
-------
Ground limestone is the only major chemical re.
quirement for this process. The calcium sulfate produced is considered to be
a waste product and is discarded. In practice, a small quantity of the basic
aluminum salt must be added periodically to make up for mechanical losses and
losses in the filter cake. Also, a tfmall amount of phosphoric acid may have
to be added to the solution occasionally to prevent the crystallization of
aluminum sulfate. The quantities of these chemicals are small and therefore
they have been classified as in-plant losses not accounted for in this process
economic study.
(4) Cost Estimate
In order to be consistent with many of the other
processes evaluated in this study, the cost of major equipment such as absorptii
columns, pumps, and blowers have been taken, whenever possible, from the
Bureau of Mines study. These costs were adjusted to 1967 purchase costs, an
were adjusted for size according to the 0. 6 power factor. Other equipment coat
59
were derived from Chemical Engineering Costs of Process Equipment files.
(a) Capital Costs
The purchased cost of major equipment for
this process, $3, 110, 000, results in an extremely high total investment of
$16, 203, 100, or $135. 03/kw capacity. See Table 53.
The largest major equipment cost is for foui
lead-lined steel absorption towers with wood-grid packing ($1, 731, 000). These
towers were sized based upon the data of Appleby for the 52 ton/day Imatra,
Finland, smelter plant which treated waste gases containing 5% SO-. The large
*•
absorbers are required because of the very small driving force which results
from the low concentration of the SO- in the gas stream, resulting in a low
equilibrium partial pressure of the SO2 in the scrubbing solution.
A large volume of the scrubber effluent
(5, 640 gprn) is heated to 100°C, stripped with a large quantity of steam, and
then cooled to 50°C for absorption. This requires the following large and
expensive heat transfer equipment:
-------
TABLE 53
BASIC ALUMINUM SULFATE PROCESS: CAPITAL COST ESTIMATE SUMMARY
ITEM FACTOR COST - $
1. Purchased Equipment 1.00 3, 110, OOP
2. Erection Labor 0.25 777, 500
3. Foundation fc Platforms 0. 18 559, 800
4. Piping 0.76 2.363,600
5. Instruments 0. 15 466, 500
6. Lisulation 0.08 248.800
7. Electrical 0. 10 311.000
8. Buildings 0.25 777.500
9. Land fc Yard Improvements 0. 13 404. 300
10. Utilities 0.40 1.244.000
11. Physical-Plant Cost 3.30 10.263.000
12. Engineering fc Construction 0. 66 2.052. 600
13. Direct Plant Cost 3.96 12.315.600
14. Contractor's Fee 0. 19 590.900
15. Contingency 0.59 1.834.900
16. Fixed Capital Cost 4.74 14.741.400
17. Working Capital. 10% 0.47 1.461. 700
Total Investment 5.21 16.203. 100
18. Capital Requirements
$/kw capacity 135.03
-------
Heat exchanger, shell and tube, stainless
steel tubes, 194, 500,000 Btu/hr duty,
194,000 sq. ft. area $683,000
Reboiler for stripping column, kettle type,
stainless steel tubes, 133, 300, 000
Btu/hr duty, 13,330 sq. ft. at
$lZ.40/ft2 165,000
Cooler for lean solution, shell and tube,
stainless steel tubes, 55,500,000
Btu/hr duty, 15,000 sq. ft. at
$8.71/ft2 131,000
Stripper overhead condenser, shell and
tube, stainless steel tubes,
78, 700,000 Btu/hr duty, 3850 sq.ft.
at$8.71/ft2 34,000
Total cost of heat exchange equipment: $1,013, 000
Other major equipment items are:
Absorption towers (4), lead lined $1,731,000
Solution pumps (9, 3 spare), centrifugal,
C.I. with SS shafts, 5, 640 gpm at
80 ft. head 38, 000
Stripping column, lead-lined steel,
6 ft. OD, 50 ft. high 39, 000
Storage tanks (6), lead-lined steel,
60, 000 gal capacity 100, 000
Filter presses (2), for scrubber effluent,
lead-lined, pressure type, 479 sq.ft. 58,000
Filter presses (2), for CaSO4, steel,
pressure type, 479 sq. ft. 29, 000
Liquefaction system, for SO- 102, OOP
Total cost of miscellaneous items: $2, 097, 000
Total major equipment cost: $3, IIO.OM
(b) Operating Costs
The estimated annual operating charges are
listed m Table 54. The total operating cost is $4, 070, 300 or $8. 56 per ton of
coal. The raw material and chemical cost for this process is relatively low. The
direct labor charge includes the labor required for the SO, liquefaction system.
b
-------
TABLE 54
BASIC ALUMINUM SULFATE PROCESS: OPERATING COST ESTIMATE SUMMARY
Fixed Capital Coat: $14, 741,400
ITEM
1. Raw Materials fc Chemicals
2. Direct Labor
3. Supervision
4. Maintenance. 5% of Fixed Capital
5. Supplies, 15% of Maintenance
6. Utilities
7. Other
8. TOTAL DIRECT COST
9. Payroll Burden, 20% of 2 It 3
10. Plant Overhead, 50% of 2, 3, 4 fc 5
11. Pack & Ship
12. Waste Disposal
13. Other
14. TOTAL INDIRECT COST
15. Depreciation, 10 % Fixed Capital/Yr
16. Taxes. 2% of Fixed Capital
17. Insurance, 1% of Fixed Capital
18. Other
19. TOTAL FIXED COST
20. TOTAL OPERATING COST
TOTAL $
91.900
120.000
43.200
737. 100
110.500
828.000
2.26
505.400
83.300
621.300
1.474.100
29. 500
14.700
1.518.300
4.070,300
2.95
1.06
18. 11
2.71
20.34
47.43
0.80
12.42
2.05
15.27
36.22
0.72
0.36
37.30
100.00
21. COST: $/Ton of Coal
22. Mill/kwh
23. BY-PRODUCT CREDIT
8.56
4.28
(SEE FIGURE 28 )
-------
The cost of utilities is $828, 000 per year.
This high cost is reflected by the large amount of steam required to strip the
SO, from the rich solution (25 moles steam per mole SO2 stripped). Also, a
large quantity of water is required to cool the lean solution and condense the
stripping steam. The utility requirements follow:
Steam: 1, 052, 000 M Ibs per year
Circulating Water: 3, 850,000 M gal per year
The CaSO, waste amounts to 14, 970 tons
per year. Adding the 5% waste rock present in the limestone and the water in
the filter cake (assumed to be 35%), the total waste tonnage is 20,830 tons.
(c) Profitability
Profitability data for the Basic Aluminum
Sulfate process are given in Figure 28 which indicates a high cost operation.
P* Ammonia-Hydrazine Process
(1) Process Description
The Ammonia-Hydrazine process involves the
scrubbing of flue gas with an aqueous solution of hydrazine salts, including
the sulfite and bisulfite. Aqueous hydrazine is added to the circulating
scrubber liquor at a rate corresponding to the rate of sulfur dioxide absorption,
and a portion of the circulating stream is continuously removed for product
recovery. The latter involves, in turn, a preliminary filtration to remove fly
ash, treatment with hydrazine to convert bisulfite to sulfite, air-oxidation of
the sulfite to sulfate, and ammonolysis of the sulfate in liquid ammonia to yield
insoluble ammonium sulfate and an ammonia-hydrazine solution. The ammonium
sulfate is separated by centrifugation, and dried by flash evaporation of the
ammonia. Anhydrous hydrazine is obtained as a relatively non-volatile liquid
by flash evaporation of ammonia, and is returned to the process. However, to
the extent that a market exists for anhydrous hydrazine a portion of the product
may be sold. In this case make-up hydrazine is supplied either through purchase
of dilute hydrazine or through the installation of a. small Raschig facility. The
Raschig method itself yields anhydrous hydrazine, but for the present purpose
only that portion of the overall facility is required which produces 3% aqueous
-------
.2
I
i
8
ol
70
60
50
40
30
20
10
-
^
•^v
Curren
v
\
^
t Priced
\
\
69 /ton
t
\
V
\
\
\
Basis : *120 megawatt power plant
• 20 million dh flue gas
• Sale of 34, 610 tons S02
per year
4567
Cost - Dollars Per Ton Coal
8
1.5 2.0 2.5 3.0 3.5
Cost - Mills Per Kilowatt Hour
4.0 4.5
BASIC ALUMINUM SULFATE PROCESS : PROFITABILITY
-------
An important distinction between thU system and
the ammonia -based systems is that only one scrubber is used in the Ammonia-
Hydrazine process. This change is probably justified because of the relatively
low vapor pressure of hydrazine as compared with ammonia; thus, hydrazine
stack losses should be negligible.
A flow diagram for the Ammonia -Hydrazine
process is shown in Figure 29.
(2) Process Reactions
Scrubber:
N2H4 + H2O + SO2 - •» N2
1/2
Mixer:
N2H5HS03 + N2H4 — * 2S0
Oxidize r:
+ 1/2
Ammoniator: ._.
(N2H5)2S04 + 2 NH3 - ^. (NH4)2S04 f + 2N;
(3) Chemical Requirements and By-Product Yields
As with all of the other processes considered in this
study, the reactant chemicals and product yields shown in Figure 29 are theoretical
quantities. Due to its low vapor pressure, hydrazine losses should be negligible;
as indicated previously, however, ammonia losses experienced in the ammonia-
based systems are significant.
The chemical requirements and by-products produced
per year for plants of three different sizes are listed in Table 55. It should be
noted that the quantity of ammonium sulfate produced is the same as that obtained
-------
A Stream
1
2
3
4
Component
S02
S02
NH3
( NH4)2S04
Lb Moles
8.36
0.42
15.9
7.94
Lbs
535
26.8
270
1048
FLUE GAS
SLUDGE AIR
SCRUBBER
MIXER
FILTER PRESS
OXIOIZER
VACUUM
EVAPORATOR
CENTRIFUGE
AMMONIATOR
FLASH TANK
AMMONIA -HYDRAZINE PROCESS : FLOW DIAGRAM
-------
TABLE 55
AMMONIA-HYDRAZINE PROCESS:
CHEMICAL REQUIREMENTS & BY-PRODUCT YIELDS
Quantity per Million Tons per Year
SCF Flue Ga8 20 Million 0. 5 Million 2. 5
Lb Mole Lb SCFH Flue Gas SCFM Flue Gas SCFM Flue!
Raw Materials
SO2 (flue gas) 7.94 508 40,230 60,350 301,730
NH3 15.88 270 21,380 32,070 160,350
By-Product
(NH4)2S04 7.94 1048 83,000 124,500 622,500
-------
(4) Cost Estimates
(a) Capital Costs
The capital cost of the Showa-Denko process
vas used as the basis for this system. One of the Showa-Denko scrubber sub-
systems, including pumps and blower, was eliminated from the present system,
resulting in a cost reduction of $240, 300. The $35, 000 dryer used in the
Showa-Denko process was also eliminated. The combined cost of the ammonia tor
ind flash tank which had to be incorporated into the system (see Table 9) were
estimated to be $35, 000. Thus, the purchased equipment cost for the Ammonia-
•Hydrazine process is $965, 300, as compared with $1, 205, 600 for the Showa-Denko
process. The capital costs are summarized in Table 56. .
(b) Operating Costs
A summary of the operating-cost estimate
is shown in Table 57 The total operating cost is $2, 765, 900, or $5. 82/ton of
coal and 2. 91 mill/kwh.
Raw material consumption is based on the
quantities shown in Table 55. In-plant losses and hydrazine make-up have not
been included in raw material usage.
Four men per shift were considered adequate
to operate the plant, a supervision requirement of one foreman per shift and one
area superintendent was assumed.
The only significant reduction, other than
that created by the lower capital costs, is the lower power requirement which is
reduced by 3, 550, 000 kwh per year due to elimination of the blower and pumps.
(c) Profitability
Figure 30 shows the profitability of the system
assuming sale of all of the ammonium sulfate produced (83, 000 tons per year).
The data indicated a cost of $3-3. 50/ton of coal if the ammonium sulfate were
sold at one-half the current selling price.
-------
TABLE 56
AMMONIA HYDRAZ1NE PROCESS: CAPITAL COST ESTIMATE SUMMARY
ITEM FACTOR COST -$
1. Purchased Equipment 1.00 965,300
2. Erection Labor 0.25 241.300
3. Foundation & Platforms 0. 18 113,800
4. Piping 0.76 733.600
5. Instruments 0. 15 144.800
6. Insulation 0.08 77.200
7. Electrical 0. 10 96.500
8. Buildings 0.25 241.300
9. Land fit Yard Improvements 0.13 125,600
iO. Utilities 0.40 386,100
11. Physical-Plant Cost 3.30 3. 185.500
12. Engineering & Construction 0.66 ^"37,100
13. Direct Plant Cost 3.96 3.822,600
14. Contractor's Fee 0. 19 183,400
15. Contingency 0.59 569.500
16. Fixed Capital Cost 4.74 4.575.500
17. Working Capital, 10% 0.47
Total Investment 5.21
18. Capital Requirements
$/kw capacity 41. 91
-------
' TABLE 57
AMMONIA-HYDRAZINE PROCESS: OPERATING COST ESTIMATE SUMMARY
Fixed Capital Cost: $4, 575, 500
ITEM
1. Raw Materials & Chemicals
2. Direct Labor
3. Supervision
4. Maintenance, 5% of Fixed Capital
5. Supplies, 15% of Maintenance
4. Utilities
7. Other
8. TOTAL DIRECT COST
TOTAL $
1,283.000
96.OOP
43.200
228.800
34,300
256,700
1,942,000
46.39
3.47
1.56
8.27
1.24
9.28
70.21
9. Payroll Burden, 20% of 2 b 3
10. Plant Overhead, 50% of 2, 3, 4 & 5
11. Pack & Ship
12. Waste Disposal
13. Other
14. TOTAL INDIRECT COST
15. Depreciation, 10 % Fixed Capital/Yr
16. Taxes, 2% of Fixed Capital
17. Insurance, 1% of Fixed Capital
18. Other
19. TOTAL FIXED COST
27,800
201,200
229,000
457,600
91.500
45,800
594,900
1.01
7.27
8.28
16.54
3.31
1.66
21.51
!0. TOTAL OPERATING COST
51. COST: $/Ton of Coal
52. Mill/kwh
BY-PRODUCT CREDIT
5.82
2.91
2,765,900
100.00
(SEE FIGURE 30)
-------
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ralue of N
M3 - $60
L.
\
Price -
tto _
in
/ton
\
1 2 3 4 5 (
Cost - Dollars Per Ton Coal
1 1 i i i i i
0 0.5 1.0 1.5 2.0 2.5 3.
Basis: 120 megawatt power plant
20 million cfh flue gas
SaIeof83,OOOtons(NHJ2S04
per year
Cost-Mills Per Kilowatt Hour
AMMONIA-HYDRAZINE PROCESS: PROFITABILITY
-------
q. Ammonia-Hydrazine Exorption Process
(1) Process Description
The Ammonia-Hydrazine Exorption process bears
the same relation to the Ammonia-Hydrazine process as does the Cominco
Exorption process to the Cominco process. Since the latter three processes
have already been discussed, no process description of the Ammonia-Hydrazine
Exorption process is considered necessary. A flow diagram of the system is
shown in Figure 31.
A variation of the Ammonia-Hydrazine Exorption
process, in which the salable produce is monohydrazine sulfate, rather than
anhydrous hydrazine and ammonium sulfate, was also considered. However,
this process has the disadvantage that monohydrazine sulfate, in contrast to
hydrazine, cannot be returned to the process if the amount produced exceeds
the demand. Further analysis showed that a single power plant producing
2.5 million SCFM of flue gas would yield more of the sulfate than the equivalent
amount of hydrazine consumed in non-military uses in this country in the year
i
1964. On this basis the process variation was eliminated from further
consideration.
(2) Process Reactions
Scrubber:
S0
2
S0
H20
N2H5HS03
2S03 + 1/2 °2 —** 2S°4
Heater:
2N2Ha
Ammoniator:
S0
H20
2 NH
(3) Chemical Requirements and By-Product Yields
Table 58 lists the theoretical raw material require-
ments and by-product yields. Reduction of the anhydrous ammonia requirement is
he major feature of this process when compared with the Ammonia-Hydrazine
process discussed previously.
-------
Reactant Chemicals Per Million SCF Flue Gas Processed
AStream
1
2
3
4
5
Component
S02
S02
so2
NH3
1 NH4)2S04
Lb Moles
8.36
0.42
6683
2.22
Lll
Lbs
535
26.8
437
38.0
146
PURIFI
oo
SLUDGE
SCRUBBER FILTER
PRESS
LIQUEFIER AMMONIATOR
FLASH TANK VACUUM CENTRIFUGE CENTRIFUGE I FLASH TANK
EVAPORATOR
GE
(NH4) S04
AMMONIA-HYDRAZINE EXORPTION PROCESS : FLOW DIAGRAM
-------
TABLE 58
AMMONIA-HYDRAZINE EXORPTION PROCESS:
CHEMICAL REQUIREMENTS & BY-PRODUCT YIELDS
Quantity per Million Tons per Year
Vlate rials
2 (flue gas)
[3
roducts
*4)2so4
2
SCF Flue
Lb Mole
7.94
2.22
1. 11
6.83
Gas
Lb
508
38
147
437
- ? fi X>f i Hi rtn
SCFH Flue Gas
40,230
3,010
11,640
34,610
OS \A1 111 rtn
SCFM Flue Gas
60,350
4,520
17,460
51,960
2R Million
SCFM Flue Gas
301,730
22,580
87,300
259,800
times 14% of SO- is oxidized to SO-.
-------
(4) Cost Estimate
The estimate for the Ammonia-Hydrazine Exorption
process will be derived from data available in the estimate for the Cominco
Exorption process.
(a) Capital Costs
The following items, included in the Cominco
Exorption process, can be eliminated:
One scrubbing tower, pumps
and blower - $240,300
Cooler in scrubber absorbent
circulating line - 53,500
Net change: - $293,800
This results in a total investment of $6, 155, 600, or $51. 30/kw capacity. The
capital cost is summarized in Table 59.
(b) Operating Costs
The total operating cost is $2,393,200, or
$5. 04/ton of coal and Z. 57 mill/kwh. See Table 60. Raw material costs were
derived from the quantities shown in Table 58.
Direct labor has been estimated at 5 men per
shift. A supervision requirement of one man per shift and an area supervisor is
considered adequate.
Utilities costs will be the same as in the
Cominco Exorption process, $756,400 minus a $32,200 reduction in the cooling
water requirement.
(c) Profitability
Two by-products are produced in this process:
liquid sulfur dioxide and ammonium sulfate. As in the case of the Ammonia-
Hydrazme process, anhydrous hydrazine could be produced for sale if desired,
depending on the existing hydrazine market.
Figure 32 provides a range of operating costs
(and profits) dependent upon the prices assigned to ammonium sulfate and liquid
sulfur dioxide. A price range of from zero to current market price is assigned to
-------
TABLE 59
ONIA-HYDRAZINE EXCEPTION PROCESS: CAPITAL COST ESTIMATE SUMMARY
ITEM FACTOR COST - $
Purchased Equipment 1.00 1, 181,500
Erection Labor 0.25 295.400
Foundation & Platforms 0. 18 212. 700
Piping 0.76 897.900
Instruments 0. 15 177.200
Insulation 0.08 94,500
Electrical 0. 10 118.200
Buildings 0.25 295.400
Land b Yard Improvements 0. 13 153, 600
Utilities Q.40 472.600
PHYSICAL-PLANT COST 3. 30 3.899.000
Engineering fe Construction 0. 66 779. 700
DIRECT PLANT COST 3. 96 4.678. 700
/
Contractor's Fee 0. 19 224.500
Contingency 0.59 697. 100
FIXED CAPITAL COST 4. 74 5.600. 300
Working Capital, 10% 0.47 555.300
TOTAL INVESTMENT 5.21 6. 155.600
Capital Requirements
$/kw capacity 51. 30
-------
TABLE 60
AMMONIA-HYDRAZINE EXORPTION PROCESS:
OPERATING COST ESTIMATE SUMMARY
ITEM
Fixed Capital Cost: $5, 600, 300
TOTAL $
1. Raw Materials & Chemicals
2. Direct Labor
3. Supervision
4. Maintenance, 5% of Fixed Capital
5. Supplies, 15% of Maintenance
6. Utilities
7. Other
8. TOTAL DIRECT COST
9. Payroll Burden, 20% of 2 & 3
10. Plant Overhead, 50% of 2, 3, 4 & 5
11. Pack «< Ship
12. Waste Disposal
13. Other
14. TOTAL INDIRECT COST
15. Depreciation, 13 % Fixed Capital/Yr
16. Taxes, 2% of Fixed Capital
17. Insurance, 1% of Fixed Capital
18. Other
19. TOTAL FIXED COST
20. TOTAL OPERATING COST
21. COST: $/Ton of Coal 5.04
22. Mill/kwh 2.57
180,600
120,000
43.200
280.000
42,000
724,200
1, 390,000
32,600
242,600
275,200
560,000
112,000
56,000
728,000
2,393.200
1.81
11.70
1.75
30.26
58.08
1.36
10.14
11.50
23.40
4.68
2.34
30.42
100.00
23. BY-PRODUCT CREDIT
(SEE FIGURE_32J_
-------
31 25 2O I5K> 5 O
00
-O
10
'5
.2"
c
v_
Q.
i_
ra
i
s
70
60
50
40
*>
20
10
0
&
Profit
-^
Loss
0123
Cost - Dollars Per Ton Coal
0.5 0 0.5 1.0 1.5
Cost - Mills Per Killowatt Hour
Basis: • 120 megawatt power plant
•20million cfh fluegas
• Sale of 11,640 tons ( NH4) S04 and
34,610 tons S02 per year 2
Current Prices :
• Liquid S02 : $69 /ton
• (MM*) SO, : $31 /ton
2.0 2.5
AMMONIA -HYDRAZINE-EXORPTION PROCESS : PROFITABILITY
-------
r. Mitsubishi Ammoniacal Liquor Process
(1) Process Description
The Mitsubishi Ammoniacal Liquor process is
markedly similar to the Showa-Denko process, and for this reason a separate
flow diagram for the process was not prepared. The two processes differ
in that the Mitsubishi process the first gas scrubber contains a more highly
concentrated ammonium sulfite -bisulfite solution, So that a vacuum evaporator
for workup of the off-stream liquor is no longer required. By cooling the liquid
effluent from the first scrubber before entry into the second scrubber, the
ammonia losses were reduced to 0. 2% of the makeup ammonia. The ammonia is
not injected into the flue gas stream as in the Showa-Denko process, but is
added directly to the scrubber liquor.
(2) Process Reactions
Scrubber:
SO2 - »• NH4HSO3
(NH4)2S03 + 1/2 02 - ~ (NH4)2S04
Mixer:
i^iouv/,, T i^m_
NH.HSO, + NH, » (NH4)2SO,
Oxidize r:
(NH4)2S03 + 1/2 02 *• (NH4)2S04
(3) Chemical Requirements and By-Product Yields
The chemical requirements and by-product yields are
the same as those shown in connection with the Showa-Denko process (see Table 8).
(4) Cost Estimate
The Showa-Denko process was utilized as the
basis for a cost estimate for the Mitsubishi process.
-------
(a) Capital Costs
The purchased equipment cost in this system
would be slightly less than in the Showa-Denko process due to the following
changes:
Eliminate evaporator -$121,000
Add cooler + 96, 000
Net change: - $25,000
This change results in a fixed capital charge of $6, 139, 200 and a total investment
of $6, 150, 900, or a cost of $51. 26 per kilowatt.
(b) Operating Costs
The operating cost would increase slightly
over the Showa-Denko system, as follows:
Lower steam requirements due
to elimination of evaporator -$6, 000
Increase cooling water +90,000
Lower fixed costs (less de-
preciation, insurance,
and taxes) -15,400
Net change: +$68,600
The increase in annual operating cost increases the unit cost per ton of coal to
$6.47 with no by-product credits.
(c) Profitability
A profitability chart was not prepared for
this system since there is only a slight variation from the Showa-Denko process.
The costs per ton of coal are as follows:
No by-product credit $6.47
Full credit for (NH4)2SO4 $1. 05
(NH4)2SO4 credit at $15. 50/ton $3. 76
-------
s. Mitsubishi Manganese Oxyhydroxide Process
(1) Process Description
In the Mitsubishi Manganese Oxyhydroxide process,
a 3 percent slurry of manganese Oxyhydroxide is utilized for the absorption of
sulfur dioxide. A portion of the absorbing slurry is continually withdrawn and
treated with ammonia and air in order to regenerate the manganese Oxyhydroxide,
which is then returned to the scrubber. The ammonia is converted to soluble
ammonium sulfate, which is removed as a 45 percent solution after clarification.
The solution is further concentrated in a crystallizer, and the precipitated crystalling
solid is separated by centrifugation and dried. A flow diagram of the Mitsubishi
Manganese Oxyhydroxide process is shown in Figure 33.
(2) Process Reactions
Scrubber:
02 » 4 MnSO4 + 2
Mixer-Oxidizer:
4 MnS04 + 8 NH3 + QZ » 4 Mn(OH)O ^ + 4
(3) Cost Estimate
Very few data for this process were found in the
literature. The flow diagram shows our conception of the system. The data
provided in the Bureau of Mines study 21 again was used for estimating purposes.
(a) Capital Costs
The fly ash scrubber subsystem is similar
to the one used as the main SO2 scrubber in the Zinc Oxide process. The purchased
equipment cost is $215, 600. The filter press was also selected from the same
system, purchased cost is $84,800.
In the Howden-I. C.I. system lime or chalk
slurry is used for removing SO2 from flue gases. Since a slurry scrubbing medium
is also used in the present process, the Howden-I. C.I. scrubber was considered
suitable. The purchased cost of this unit is $263, 600.
-------
Reactant Chemicals Per Million SCF Flue Gas Processed
FLUE GAS
PURIFIED GAS
Astream
1
2
3
4
Component
soz
S°2
NH3
( NtU SO.
a2 *
Lb Moles
8.36
a 42
15.9
7.94
Lbs
535
26.9
270
1048
FLY ASH
FLY ASH SCRUBBER
FILTER
S02 SCRU68ER
MIXER-0X1DIZER
CLARIFIER CRYSTALLIZER DRYER
CENTRIFUGE
MITSUBISHI MANGANESE OXYHYDROXIDE PROCESS : FLOW DIAGRAM
-------
Two 150, 000 »gal lead-lined steel tanks
similar to those specified in the Magnesium Hydroxide process were selected
for the air oxidizer and mixer subsystem. The purchased cost, including an
air compressor, is $93,400.
The clarifier in the Zinc Oxide system was
selected for use in that step in which the magnesium oxyhydroxide is allowed to
settle. The purchased cost is $46, 800.
A large amount of ammonia is used in this
process. A storage system similar to that specified in the Fulham-Simon-Carves
process was selected. The equipment required is as follows:
Pressure sphere, 229, 000 gal, with
refrigerator and compressor $105, 500
Storage tank, 3,935 gal, for 15%
NH3 liquor 7, 200
Head tank, 3,935 gal, for 15%
NH3 liquor 7,200
Total: $119,900
Although an evaporator is not needed in
this process, a crystallizer is required. Cost File 133 lists complete
costs for crystallizer and dryer systems for ammonium sulfate, which include
the vacuum system, instrumentation, pumps, centrifugation and drying equipment,
insulation and piping. In terms of purchased cost, this amounts to $600 per ton
of ammonium sulfate per day. At 250 tons per day, the purchased equipment
cost is $150,000.
The cost of various pumps was estimated
at $41,400.
The total purchased equipment cost is
$1, 015, 500. The total investment is $5, 290, 800, or $44. 09/kw capacity. See
Table 61.
-------
TABLE 61
MITSUBISHI MANGANESE OXYHYDROXIDE PROCESS:
CAPITAL COST ESTIMATE SUMMARY
ITEM FACTOR COST - $
Purchased Equipment 1.00 1.015. 500
Erection Labor 0.25 253. 900
Foundation & Platforms 0. 18 182. 800
Piping 0.76 771,800
Instruments 0. 15 152. 300
Insulation 0.08 8 1 . 200
Electrical 0. 10 101. 6QQ
Buildings 0.25 253. 900
Land & Yard Improvements 0. 13 132.000
Utilities 0.40 406.200
Physical -Plant Cost 3.30 3.351.200
Engineering It Construction 0. 66
Direct Plant Cost 3. 96 4.021.400
Contractor's Fee 0. 19 192. 9QQ
Contingency 0.59 599. 2QQ
Fixed Capital Cost 4. 74 4.813.5QQ
Working Capital, 10% °-47 477. 3QQ
Total Investment 5.21 5.2QQ.8QQ
Capital Requirements
$/kw capacity 44. 09
-------
(b) Operating Costs
The raw material requirements for the
Mitsubishi Manganese Oxyhydroxide process are shown in Table 62 and the
operating costs in Table 63. The latter totals $2, 849, 300, or $6. 00/ton of
coal and 3. 00 mill/kwh.
Direct labor was estimated at 5 men per
shift with one foreman per shift, and one area superintendent.
Since dependable data were not available,
the utilities requirements wore assumed to be the same as those of a similar
process, the Magnesium Hydroxide process.
(c) Profitability
Figure 34 shows the profitability for the
Mitsubishi Manganese Oxyhydroxide process. It appeared that the net cost
would be $3-4/ton of coal and 1. 5-2. 0 mill/kwh.
-------
TABLE 62
MITSUBISHI MANGANESE OXYHYDROXIDE PROCESS:
CHEMICAL REQUIREMENTS & BY-PRODUCT YIELDS
Quantity per Million Tons per Year
SCF Flue Gas 2Q Mlllion 0 5 Mlllion 2. 5 Million
Lb Mole Lb SCFH Flue Gas SCFM Flue Gas SCFM Flue Gas
Materials
i, (flue gas) 7-94 508 40,230 60,350 301,730
ft
I3 15.88 270 21,380 32,070 160,350
roduct
H4)2S04 7.94 1048 83,000 124,500 622,500
-------
TABLE 63
MITSUBISHI MANGANESE OXYHYDROXIDE PROCESS:
OPERATING COST ESTIMATE SUMMARY
Fixed Capital Cost: $4, 813, 500
ITEM
1. Raw Materials b Chemicals
Z. Direct Labor
3. Supervision
4. Maintenance, 5% of Fixed Capital
5. Supplies, 15% of Maintenance
6, Utilities
7. Other
8. TOTAL DIRECT COST
9. Payroll Burden, 20% of 2 & 3
10. Plant Overhead, 50% of 2, 3, 4 & 5
11. Pack fa Ship
12. Waste Disposal
13. Other
14. TOTAL INDIRECT COST
15. Depreciation, 10 % Fixed Capital/Yr
16. Taxes, 2% of Fixed Capital
17. Insurance, 1% of Fixed Capital
18. Other
19. TOTAL FDCED COST
20. TOTAL OPERATING COST
21. COST: $/TonofCoal
22. . Mill/kwh
6.00
3.00
TOTAL $
1,282,800
120,000
43, 200
240,700
36,100
248,100
1,970,900
32,600
220,000
252,600
481,400
96,300
48,100
625.800
2,849,300
69.17
23. BY-PRODUCT CREDIT
(SEE FIGURE 34) _
-------
1.61
~ 1-15
'E
OL. « «£
TO
I .69
o>
£ .46
a.
.23
CM
25
fe 20
a,
5 15
B
8 10
i_
a.
^
Currer
\
\
it price
\
-$31/
ton
Maximum price
(Equivalent to valu
x
N
\
eof ~
n
Basis : *120 megawatt power plant
•20 million cfh flue gas
•Sale of 83,000 tons (NHd) S04 per year
1234567
Cost - Dollars Per Ton Coal
0.5 1.0 1.5 2.0 2.5 3.0 3.5
Cost - Mills Per Kilowatt Hour
MITSUBISHI MANGANESE OXYHYDROXIDE PROCESS : PROFITABILITY
-------
t. Mitsubishi Lime Process
(I) Process Description
A 10% lime or limestone slurry is used in the
Mitsubishi Lime process as the absorbent for sulfur dioxide. The process is
similar to the Howden-I. C. I. (Cyclic Lime) method. The main difference is
that calcium sulfate of high purity ).s obtained in the Mitsubishi method; it will
be recalled that the calcium sulfate produced in the Howden process is a
highly contaminated, n on salable material. The high purity of the calcium
sulfate product is presumably a consequence of working with a scrubber
effluent solution which is free of fly ash.
(2) Process Reactions
< Scrubber:
S02 + Ca(HC03)2 — » CaS03
CaSO3 + 1/2 O2 — »• CaSO4
Makeup Tank:
CaO + CO - •» CaCO
3
and/or
CaCO^ 4 + H,O + CO," •» Ca(HCO-),
•J * u C, 36
Oxidize? Tank-
CaSO3 | + 1/2 O2 —•» CaSO4 J
(2) Chemical Begutrements and Waste Product Yields
Table 64 gvves the raw material requirements and the
procucts for creating one million SCF of flue gaa by the Mitsubishi process.
Although ihe data Khotm are applicable for the case where lime (calcium oxide) is
used, slaked June (caie.ium Wdroxide) or limestone (calcium carbonate) can be
employed
(4) Cost Estimate
Since this system ia oimilar to the Howden1-!. C. I.
prc .ess, the UlUi -«^ u^j ;.„. the basis for the cost estimate.
-------
TABLE 64
MITSUBISHI LIME PROCESS:
CHEMICAL REQUIREMENTS & BY-PRODUCT YIELDS
FOR THE SYSTEM USING LIME
Quantity per Million ^ Tons per Year
SCF Flue Gas 20 Million 0. 5 Million 2. 5 Million
Lb Mole Lb SCFH Flue Gas SCFM Flue Gas SCFM Flue Gas
iw Materials
S02 (flue gas) 7.94 508 40,235 60,350 301,730
CaO 7.94 445 35,215 52,825 264,115
'-Product
CaSO4 7.94 1367 108,300 162,450 812,250
-------
(a) Capital Costs
The reference describing the Mitsubishi wet
processes .states that the equipment cost of the Lime prpcess is about twice that
of the Sirrplified Lime processes. As indicated previously, the Simplified
Lime prot ess is considered identical to the Howden-I. Ct I. process. On this
basis, the capital cost for the Lime process was obtained by doubling the
Howden-I C.I. capital cost, resulting in a fixed capital cost of $4,839, 500 and
a total in\ estment of $5, 319, 400. This is equivalent to $44. 33 per kw capacity.
(b) Operating Costs
*" " ' """ 'V " The operating coai'summary'is'shown iri"~ *"'
Table 65, which shows an annual cost of $1, 981, 900, or $4. 17/ton of coal and
2. 09 mill 'kwh.
Raw materials charges, remain the same
as in the Iowden-I. C. I. process.
Direct labor was increased from four^to
five mien ier shift due to the addition of the drying system.
iJtilit'es charges were increased to handle ,
the additi -nai operations needed to obtain pure calcium sulfate ,
•s. i n. ,
Waste disposal costs were eliminated since
ttie pro e '.i" ic-f v^r-fca ;o provide hi^h-pi-rUy f >1 -u»rn sulfate.
ir t-.ijr t>Tocess the calcium sulfate is produced
.-tiu-.Abl-/ b? soJci. * Tigt re 35 shovs tbe .vrofita-
various credits to this by-product.
* „.
Tae Ph.;se Devaluation, showed tJiat the market would be limited for this material
in the I mted States. " " '' "" ** "
-------
TABLE 65
MITSUBISHI LIME PROCESS: OPERATING COST ESTIMATE SUMMARY
Fixed Capital Cost: $4, 839, 500
ITEM
1. Raw Materials & Chemicals
2. Direct Labor
3. Supervision
4. Maintenance, 5% of Fixed Capital
5. Supplies, 15% of Maintenance
6. Utilities
7. Other
8. TOTAL DIRECT COST
9. Payroll Burden, 20% of 2 b 3
10. Plant Overhead, 50% of 2, 3, 4 & 5
11. Pack & Ship
12. Waste Disposal
13. Other
14. TOTAL INDIRECT COST
15. Depreciation, 10 % Fixed Capital/Yr
16. Taxes, 2% of Fixed Capital
17. Insurance, 1% of Fixed Capital
18. Other
19. TOTAL FIXED COST
20. TOTAL OPERATING COST
21. COST: $/Ton of Coal
22. Mill/kwh
4. 17
2.09
TOTAL $
545,800
120,000
43,200
242,000
36.300
112.000
1,099.300
32.600
220.800
253,400
484.000
96.800
48,400
629.200
1.981.900
27,54
6.06
2. 18
12.21
1.83
5.65
55.47
1.64
11. 14
12.78
24.42
4.88
2.44
31.75
100.00
23. BY-PRODUCT CREDIT
(SEE FIGURE 35 )
-------
u. Mitsubishi Red Mud Process
M Process Description
The absorbent used in the Mitsubishi Red Mud
process is a slurry of red mud, the residue resulting from the extraction of
alumina from bauxite. In the simplest application of this method, the scrubber
effluent is discarded after sulfur dioxide absorption. This is the approach on
which the present analysis is based. In another version of this process, one
for which very little information is available, the solid residue remaining
after absorption is separated into the fractions which are rich in alumina,
silica, and ferric oxide, respectively. These are utilized as industrial
materials.
(2) Process Reactions
The component in red mud which reacts with SO-
may be considered to be sodium oxide.
Scrubber:
+ 1 12 O2 - »•
(3) Chemical Requirements and By-Product Yields
Table 66 gives the raw material requirements and
waste -product yields for the process. The waste products are shown as dry
weights; in addition, it was assumed that approximately 35% free water will be
included in the total weight.
(4) Cost Estimate
Due to the very limited data on the Red Mud process,
the estimate was made only on the basis of a once-through operation, i.e. , the
spent red mud is simply discarded after use.
(a) Capital Costs
It was assumed that the Howden-I. C. I. process
equipment could be adapted to this system. Accordingly, the same fixed capital
-------
TABLE 66
MITSUBISHI RED MUD PROCESS:
CHEMICAL REQUIREMENTS & BY-PRODUCT YIELDS
Quantity per Million Tons per Year
—SCP Flue Gas 20 Million 0. 5 Million 2. 5,Million
Lb Mole Lb SCFH Flue Gas SCFM Flue Qas SCFM Flu'« Cat
Raw Mats rialH
S02 (flu* gas) 7.94 508 40,130 60,350 .301,7.30
Red Mud (dry) 7.94a ^6000b 475,200 712,800 3,564,000
By-Product
Red Mud Slurry - 6508 515,430 773,150 3,865,730
containing re-
acted 3
-------
cost of $2,419, 800, or a capital requirement of $22. 16 per kilowatt capacity,
was assigned to the Red Mud process.
(b) Operating Costs
Where applicable, the various cost elements
of the Howden-I. C. I. process were used.
It was assumed that a plant using red mud
slurry for removal of SO, from flue gas would be located near the bauxite
processing facility. No charge was made for acquisition of the raw material.
The waste disposal costs are very high due
to the quantity of waste products which have to be handled.
Table 67 summarizes these costs. The
$7. 58 per ton of coal operating cost would be reduced to $1. 72 per ton if there
was no cost associated with waste disposal (this assumes that the adjacent
bauxite plant which generated the red mud has suitable disposal facilities on-site).
(c) Profitability
The profitability of this system is not
indicated in graphical form. Its application as a process is limited by the
availability of red mud since one 1400 megawatt power plant (Case III) would
generate enough SO, to consume approximately all of the red mud available in
£
the United States. It is doubtful that the process would be economical if disposal
of waste products represents a substantial cost. Adequate data were not available
to evaluate conversion of these waste products into salable industrial materials.
-------
Y A3 Jut: 67
MITSUBISHI RED MUD PROCESS: OPERATING COST ESTIMATE SUMMARY
Fixed Capital Cost: $2,419, 000
ITEM
1. Raw Materials fc Chemicals
2. Direct Labor
3. Supervision
4. Maintenance, 5% of Fixed Capital
5. Supplies, 15% of Maintenance
6. Utilities
7. Other
8. TOTAL DIRECT COST
9.
10.
11.
12.
13.
14,
15.
16.
17,
IS.
'/A.
22.
Payroll Burden, 20% of 2 & 3
Plant Overhead, 50% of 2, 3, 4 & 5
Pack fit Ship
Waste Disposal
Other
TOTAL INDIRECT COST
Depreciation, _ IQ % Fixed Capital/Yr
Thxe-*, 2% ui' Fixed Capital
Insurance, 1% of Fixed Capital
Othe?
TOTAL FIXED COST
10TAL OPERATING COST
COST- $'TraofCoal 7.58
Mi)l/kwh 3. 79
TOTAL $
-0-
96.OOP
43. 200
121.000
18.200
56.200
3S4> 600
27,800
139,200
2.783.300
2.950.300
242.000
48,400
24.200
314,600
3^99500
2.67,
1.20
o.si,'1
• 1. 56 fr
0. 77
3.87
77.32
i
6.72
1.35,
0.67'
8.74
100.00.
2b. BY-PRODUCT CREDIT
NONE
-------
v. Other Processes
Several of the processes listed in Table 1 were evaluated
as part of the Phase I effort.. Many of these are under current development in
this country and are, in large part, proprietary, so that available data are not
sufficient to permit a meaningful analysis. For such processes, only the individual
contractors involved in the development of these processes can adequately provide
economic analyses of the type considered in the preceding sections of the report.
By mutual agreement with NAPCA for the present program,
the evaluation of the following processes was considered to lie beyond the scope
of the present effort:
• Wisconsin Electric Power
• Wisconsin Electric Power/Universal Oil Products
• Combustion Engineering
• Bechtel
• Ionics/Stone and Webster
Some information was supplied by NAPCA relating to
the Wellman JLord/Beckwell process, and an evaluation was planned for this
process. However, a large portion of the process data were undisclosed in
the literature provided, and it was finally concluded that a meaningful analysis
569
could not be carried out. A Belgian patent, which became available at a
later date, provided a general description of the process. Potassium sulfite in
an aqueous system is used as the absorbent to remove the SO- from the flue gas.
Three major operations are used in the process: reactor, crystallization, and
stripping. In the first step, a part of the excess potassium sulfite and the SO-
react with the formation of potassium bisulfite. In the second operation, the
solution is cooled resulting in the transformation of potassium bisulfite to potassium
pyrosulfite during crystallization. After filtration the crystals of potassium
pyrosulfite are heated in the third operation yielding potassium sulfite and SO-
£ •
The potassium sulfite is returned to the reactor operation while the SO, is
recovered as liquid SO- or as sulfur. The crystallization of potassium pyro-
C*
sulfite and its conversion to potassium sulfite and SO apparently are the main
economic advantages of the process. Potassium sulfate formed, due to the
presence of SO , has to be removed periodically. In addition, the potassium
-------
sulfite may react with the oxygen in the flue gas to yield more potassium sulfate.
It may be necessary, therefore, to use oxidation inhibitors such as hydroquinone
to prevent this reaction. A laboratory program comparing the behavior of potaeaiun
and sodium ions as absorbents was of interest and is reported in Volume II of thit
Final Report.
Of the entries in Table 1 only the Kanagawa, Guggenheim,
and Diethylenetriamine processes, remain to be discussed. Of these the Kanagawa
process could not be evaluated because of the sparcity of available data. However,
it was evident from the single reference available on this process that the
scrubber medium is a naturally occurring water, and consequently it will exhibt ,
a low capacity for SO,. On this basis, the process may be compared with the
Haenisch-Schroeder, Battersea, and other processes involving a water scrubber
and, as will be seen in a subsequent section of the report dealing with process
selection (Section II, C, 3), this type of process is among the poorest candidates
for a continuing effort in Phase III of the present program. In particular, the
capital cost of all such processes is exorbitant, because of the large scrubbing
towers needed to offset the low absorbing capacity of the scrubber medium.
The Guggenheim process was not evaluated because it
differs substantially from the Cominco Exorption process only in that in the
former process recovered SO- is reduced to elemental sulfur by heating with
coke. This step is applicable in general whenever SO, appears as a product,,
ana the conversion of SC, to other salable forms of sxilfur, including elemental
sul"\:r, viis Lus t.c*ced a& a e?paiatc& aectiju of this report (see Section III. C. 3.e).
Aqueous solutions of various aliphatic amines have been
P*- )j,'ossd from time to time ai substitutes for aqueous ammonia in the Cominco
f_> jrptif r prc^^r. The main advantag- oi an amine over ammonia is its
i'jtn*i*fi -.or -,oiatiUty, oo that during the thermal regeneration of SO-, little or
It
no amint 33 losf, and the regeneration step can be carried to any desired degree
ci" coir7 etion, 1 -nue^ only by the cost of steam. Diethylenetriamine is one
example of ? pmoosed amine, The Diethylenetriamine process was not fully
evaluated on the curient program, btci-use only initial laboratory studies had
been reported, so that very few data were available on which to base an economic
analysis. The tollowing r*marko, however, are relevant to this process in
particular, and to the u»c A tunes in general.
-------
Aliphatic amines are more basic than ammonia
= 1.8 x 10 ) so that absorbed SO0 is more tightly held. This situation is
^ 224
unfavorable to thermal regeneration and, in fact, Johnstone has indicated that
for overall minimum steam requirements the absorbing amine should exhibit a
—ft
base constant of about 10 . Another general limitation to the use of amines is
the limited solubility of amine sulfites in water relative to that of ammonium
sulfite. This is particularly true of amines containing a relatively high C/N
ratio, in which case the effect is a decrease in the capacity of the absorbing
solution.
The polyethylene amines, of general formula
H,N(CH,CH7NH) CH^CH-jNH,, where x < 4, contain secondary amine functions
£• £* £* X Lt {* t» —^~
whenever x > 1. In general, secondary amines exhibit larger base constants
tha.n primary amines, so that if x > 1 the polyamine should be more basic than
et,iylenediamme (x = 1, KR = 8.5 x 10" ).
The ethanolamines, H N(CH9CH9OH) , where x >0 and
x c* £• y
\ - y = 3, were found by Johnstone to be readily oxidized in experiments in-
volving SO_ absorption, even in the absence of air, and thus appear to be
eliminated as scrubber components on this basis alone.
3. Process Selection
a. Introduction
A comparative economic analysis of the processes in this
study was needed in order to select candidates for Phase III. Several important
factors which affected this assessment include.
• Total capital investment and investment per
kilowatt of generating capacity (see Figure 36)
• Operating cost expressed as dollars per ton of
coal consumed or as mills per kilowatt hour
(see Figure 37)
o By-product utilization and/or disposal, and
justification of credits applied to operating
cost.
• The impact of the by-product on the economy
of the United States.
-------
Figure 36
C'V.PARr'.YiVE ASSESSMENT OF AQUEOUS 8ASE& PROCESSES FOR
fcEKdVING S02 FROM FLUE GASES
CAPITAL INVESTMENT
s> R o
MILLIONS OF DOLLARS
? 4 6 8 10 12 14 1£ 18 20
26 28 30 32 34 3S
5.;?-' -r>; - -i 1
^ - .. -a» -roj^ -t«si
OXYHYWC: IK
MAGKESiUM HYDROXIDE
OZO^-MniON (36SEC)
BASIC ALUMINUM SULFATE
HAENiSCH - SCHROEDER
0 10 20 3040 50 60 70 80 90 100 110 120 130 140 150 160 170 180
-------
Figure 37
COMPARATIVE ASSESSMENT OF AQUEOUS BASED PROCESSES FOR
REMOVING S02 FROM FLUE GASES
OPERATING COSTS
r
DOLLARS/TON OF COAL
-COST—f- PROFIT -*-f
PROCESS
14 13 12 11 10 9 8 7 6543
BATTERSEA
ZINC OXIDE
AMMONIA-HYDRAZINE EXORPTION
AMMONIA-HYDRAZINE
MITSUBISHI MANGANESE OXYHYDROXIDE
MANGANESE OXIDE
MITSUBISHI LIME
SHOWA - DENKO
MITSUBISHI AMMONIACAL LIQUOR
COMINCO EXORPTION
FULH AM -SIMON -CARVES
MAGNESIUM HYDROXIDE
HOWDEN I.C.I. & MITSUBISHI SIMPLIFIED LIME
COMINCO
WET THIOGEN
MAGNESIUM OXIDE "
BASIC ALUMINUM SULFATE
MITSUBISHI RED MUD
OZONE - Mn ION
SULFIDINE
HAENISCH - SCHROEDER
I — '
. * * * *
BY-PRODUCT CREDIT
" 1 T
0%
a A
5>
jr
s
C
b. t
a 10
NTICIPATED. AS FC
Dz - $20/TON (CURR
>IH ) SO. - $15.50/T
* Z *
- $45/TON(LONG)
aS04 - $3.68/TON ($
TO SALABLE BY-PR
1
1
1 L__
1
0% b
LLOWS
LNT PRICE $69)
ON ($31)
($38)
4.25)
ODUCT
1
1
1
1
1
^
c
1
ll_
1
C
4EUTRAL W/
1
1
1
1
TER-I
1
1
1
.3=3
I
1
1
1
r
1
1
1
1
i
<
1
1
1 ? f
II I
|«- THAMES RIVER WATER
1
m
i
i
i
i
i
i
i
i
i
ts)
L-31 30 29 28 27 26
By-product wHur fuNy credited
Aaumw 2856 oxidation of removed. SO2 to MgSO4 wWeh for
10* oxidation WM
-------
Consideration of capital investment and operating
this analysis was based on comparative and not on absolute values due to the natuti
of the cost estimates used. Byproduct generation could be analyzed in a more
absolute fashion since the annual quantity of SOg emissions from power plants was
easily related to the amounts of the various by-products obtained. The value
applied to these by-products, as mentioned before, was critical. The range of
by-product credit shown for each process in Figure 37 included extreme cases in
which either no credit or full credit was allowed, based on prices which were
current in February 1968. Neither of these cases was considered realistic.
Within this range was included an anticipated credit, based on anticipated by-
product selling prices.
Of the sulfur-containing products derivable from the variou.
processes evaluated, ammonium sulfate appeared to be the least desirable, at leasi
in this country, for the following reasons:
• Ammonia, ammonium nitrate, and urea, all of
which contain a much higher percentage of
nitrogen than that in ammonium sulfate, are
readily available as fertilizers.
« The sulfate radical contributes very little to
the nutrient value of ammonium sulfate when
this compound is used as a fertilizer.
w Th'i rerroval o£ suii'ur, if the form Oi
ammonium soliat'., from the flue gas
derived from only three 1400 megawatt
ponvtr plants would supply the present
market for ammortom sulfate in this
country
For the above reasons it was decided to eliminate, as a candidate for further study
in Phase III, any process which affords ammonium sulfate as the major salable
product.
-------
b. Processes Eliminated
Although high capital investment was not considered as the
only criterion for rejection of a specific process, the high investment is reflected
in the operating cost by the contribution of depreciation and other fixed charges.
High capital investment and operating costs accounted for the elimination of
several processes:
• Haemsch-Schroeder - this process is expensive
because of the high cost of stripping SO- from
large quantities of water.
• Ozone-Manganese Ion - a. large, expensive
scrubber is needed to compensate for the low
absorption rate and capacity. The dilute
sulfuric acid generated as a by-product has
little utility.
• Basic Aluminum Sulfate, Sulfichne, Wet
Thiogen - low capacity in each case results
in the need for a large, expensive scrubbing
system.
• Mitsubishi Red Mud - although the capital cost
is low for this process, the operating cost is
high because of disposal of the spent mud, In
addition, the amount of red mud available would
limit the process to the treatment of flue gas
from a single 1400 megawatt plant.
The processes using ammonia as the absorbing medium
and which generate ammonium sulfate as the only or major by-product were con-
sidered eliminated for the reasons considered in the preceding section. They
all require high credits for ammonium sulfate to be attractive. The following
processes are included in this group:
• Fulham-Simon-Carves
• Showa-Denko
» Cominco
• Ammonia Hydrazine
• Mitsubishi Manganese Oxyhydroxide
• Mitsubishi Ammoniacal Liquor
• Magnesium Hydroxide
-------
A few processes have the disadvantage that the SO., is not
recovered in a usable form. The Battersea, Howden-I.C.I., and the Mitsubishi
Red Mud and Simplified Lime processes are typical examples. , They have the
additional disadvantage that the by-products generated cause equivalent contami-
nation in solid form. In the Battersea process, the calcium sulfate produced is
discharged into the water source, whereas the filter cake from the Howden-I.C.I,
process must be dumped on land or in the sea. In isolated cases, however,
where the cost of waste disposal can be substantially reduced, and where this
waste would not create a serious solids pollution problem one or more of these
processes might be of interest, especially for existing power plants.
c. Marginal Processes
With the elimination of the various processes considered
above, only six processes remained for consideration. Three of these are of the
same type, and are considered in this section. These include the Zinc Oxide,
Magnesium Oxide, and Manganese Oxide processes.
All of the processes of interest are regenerative in the
sense that absorbed SO, is recovered by calcination of a metal salt. A valueless
by-product is produced in each case because of oxidation of the SO,.
594
Johnstone considered the use of magnesium oxide in his
development of the Zinc Ox?de process, and concluded that it is much less desirable
on tr*> botsis tha+ the calcination temperature is considerably higher for the
magnesium compound, and that the ^-?coi"ipos^tion leads to products other than
C»O2, tV. ;j3incval One b^v4?^ *u3iur trtoxide-. Thu&, zinc sulfite, at 500°C, yields
•zJzic oxide and pure jiOg, whereas magnesium sulfite must be heated to 1000°C, and
is part7, / c..nve,.;< d z»+ this temperature «••> valueless magnesium sulfate. The liter-
ature sv-veyed indicated that loss of SO- through oxidation in the Magnesium
Oxide process ta about twice that in the Zinc Oxide process.
The above arguments were considered to justify the choice
01 the Z^c v>«idj process over the Magnesium Oxide process. Additional factors
favoring ihe Zmc Oxide process are its much lower operating cost, slightly lower
capital investmen% and the fact that calcium, rather than the more expensive
magnesium, is discarded, as v.?ste.
-------
The Manganese Oxide process suffers from many of the
deficiencies considered above for the Magnesium Oxide process. Although
available information was not considered adequate to permit an unequivocal
evaluation of the process, it is known that calcination requires temperatures
in the range 1000 to 1100 C, and sulfur trioxide appears as a decomposition
product in the form of a dilute and valueless mist. Moreover, the process
involves an oxidation-reduction reaction, rather than neutralization, in the
scrubber, and this results in a relatively long residence time. As in the case
of the Magnesium Oxide process, a comparison of the Manganese Oxide process
with the Zinc Oxide process showed unfavorable capital and operating costs.
The conclusion which was drawn from the above discussion
was that the Zinc Oxide process was much to be preferred as compared to any of the
available processes in which SO. is regenerated through calcination. Accordingly,
the Zinc Oxide process was chosen as a candidate for Phase III, and the Magnesium
Oxide and Manganese Oxide processes were rejected.
d. Candidate Processes
With the elimination of the Magnesium Oxide and Manganese
Oxide processes, four processes remain, and each of these was considered to be
a candidate for continuing effort in Phases II and III. The candidate processes,
which will be discussed in turn, are the following:
• Zinc Oxide
» Ammonia-Hydrazine Exporption
• Commco Exorption
• Mitsubishi Lime
(1) Zinc Oxide Process
Although no recent work has been done on the Zinc
Oxide process, the analysis indicated a relatively low capital investment and one
of the lowest operating costs for the process. It also has the advantage of
generating a desirable by-product
-------
A disadvantage of the process is the multi-unit
operations needed in the regeneration of the absorbent and in the recovery of
SO2, which in turn implies the need for a relatively large area in which to house
the process. Although calcium su.Lfate is produced in small quantity, it can be
considered to represent a further deficiency of the process.
Phase II of the program, which encompasses the
overall laboratory effort, was devoted in part to the improvement of the Zinc
Oxide process. Areas which were considered of immediate interest included
lowering of the temperature required for calcination, preferably to the point of
elimination of the calciner so that SO, can be recovered through the use of
steam; and, elimination of, or at least a significant reduction in , the extent of
oxidation, so that little or no calcium sulfate appears as waste. To the extent
that oxidation could be minimized, a corresponding simplification in process
citjuipment would be attainable.
Because extensive data were available pertaining
to the development of the Zinc Oxide process, the effort on Phase III of the
program was concerned initially with tnis process. Process improvement
resulting from the laboratory program in Phase II was to be incorporated in
the Phase III effort. The brief Phase II program did not provide any process
improvements.
{2} Ammonia-Hydr? sine Exorption Process
The economic analysis indicated reasonably low
capital and operating costs for tins process. The main product is SO,, as in
the Z?nc Oxide process, but ;he by-product in the present case is ammonium
s \Uatc*, si, that nx wa bte product requiring disposal is obtained. To the extent
'•hat ann\d_"-Ai? Lydraaine could be sold, tins product would also be available
/•"cm tae process on A stand-by basis.
If no hydrazine were sold, no hydrazine make-up
would be requir d except tl^i due to liquid leakage in the system. Losses due
to volatility were expected to be small, inasmuch as Johnstone has shown224 that
for sulfite-bisulfita scrubbing media in the temperature range 35 to 90°C even
metft./iarmne fbcilmg point -7°C) is far superior to ammonia (-33°C) in this
-"sinect. Thus, onthiib-v s * scrubber medium containing hydrazine (+114°)
-------
The economic analysis for the Ammonia-Hydrazine
Exorption process did not include a cost for the synthesis of aqueous hydrazine.
This material could be purchased as 85% hydrate for startup and, i£ no
anhydrous hydrazine were sold, as makeup for small plant losses. However, plant
makeup could also be supplied in the form of 3% liquor, which could be produced at
relatively low cost. Since the latter contains 97% water, its transport to the
plant site would be uneconomical, an alternative route would involve the in-
corporation of a hydrazine miniplant into the flue gas scrubbing complex of
sufficient size to just provide for the loss incurred. Although the cost of such
a plant was not considered in the Phase I effort, it would be of sufficiently
small size so that its effect on the capital and operating costs for the overall
process would be nearly negligible. Such a plant could also be considered for
use in providing material for the plant startup.
If hydrazine is offered for sale, the size of the
miniplant would be increased in a commensurate manner, unless an over-
capacity had been allowed originally. The effect on costs would then be more
significant, but would be offset by the credit obtained for the hydrazine sold,
and in the most favorable case, the overall flue-gas scrubbing facility could be
operated at a profit.
The use of hydrazine sulfite-bisulfite as a scrubber
medium appeared highly desirable when compared with other sulfite-bisulfite
systems, from the standpoint of capacity of the scrubber for SO?. Thus,
224
Johnstone noted that although sodium sulfite-bisulfite had the advantage of
non-volatility, the system capacity was severely limited by the solubility of the
salts (approximately 8 moles/100 moles of water). The analogous ammonia
system showed good capacity (22 moles/100 moles), but was limited by the
volatility of ammonia. The latter not only affected scrubber losses, but im-
plied a limit in the extent to which SO, could be regenerated with steam from
the ammonia scrubber effluent. In the case of hydrazine, volatility should be
negligible. At the same time the melting points of hydrazine salts je. g. , di-
hydrazine sulfite, m. p. 6l°C, dihydrazine sulfate, m. p. 87 (compare
™"D
ammonium sulfate, m. p. 513° with decompositions are such that it appeared that
hydrazine sulfite-bisulfite mixtures would exhibit melting points below the normally
employed scrubbing temperature for flue gases of 50 C. Thus, the scrubber
medium could require no water, although an equilibrium involving a small amount
-------
of water would be reached because of the water content of the incoming flue gas,
and that introduced in connection with make-up hydrazine. The overall affect
would oe that of dramatically increasing the capacity of the scrubber for SO2,
inasmuch as water, an inert solvent with Lttle absorbing capacity at 50 for
SO2, for the most part need not be cycled. It may also be noted, in this con-
nection, that molecular oxygen is much less soluble in concentrated salt
solutions, so that as the composition of the scrubber medium approaches that
of a fused salt mixture the extent of oxidation in the scrubber should be corres-
pondingly reduced.
With respect to the desorption of SO- on heating
224
the rich scrubber liquor, Johnstone has shown that the steam requirements
can be directly related to the lonization constant of the base from which the scrubber
liquor is formulated. The ideal base constant has been determined to be approxi-
8 -10
mately 10. Weaker bases (such as xylidme, Kfi - 10 ) are good desorbers but
poor absorbers, whereas the reverse is true for stronger bases. On this basis,
hydrazine approaches the desired base strength (K— = 8. 5 x 10 ) more closely
than any of the other bases which have been considered for flue gas scrubbing
(NH3, Kfi = 1. 8 x 10"5; CH3NH2, K^ = 4. 4 x 10~4), and consequently it was
believed that the steam requirements for the liberation of SO, from hydrazine-
based systems would be minimal.
Part of the Phase II effort was devoted to a demon-
stration of process feasibility. Although hydrazine was found to be an excellent
absorber for SO,, attempts to regenerate the hydrasine for reuse resulted in its
being oxidired (see Part Three). Consequently, no work was done on Phase III
for the Ammonia-Hydrazine Exorption process.
(3) Cominco Eocorption Process
The Cominco Exorption process involves somewhat
hxgher capital a,-d operating costs than those for the other candidate processes.
Steam costs and the consequent" need for huat exchangers contribute heavily to the
overall cost., ^'he process \s similar to the Ammonia-Hydrazine Exorption
process in that SO? and ammonium sulfate are produced as salable products, and
no waste is formed. Although the procrss has been used chiefly in smelter appli-
cations, xt has also been cor,Moered for use with flue gases containing 0. 3% SC^.
O; all the ammonia processes considered in Phase I, only the Cominco Exorption
process yields SO2 as the major sulfur-containing product.
-------
Although the two exorption processes considered
in this section are of the same type, it was not considered desirable to eliminate
either as a candidate process. The Ammonia-Hydrazine Exorption process should
exhibit several unique features, as discussed above. However, the process had
not been demonstrated, and consequently could not logically be compared to the
established Cominco Exorption process. The latter, in turn, would appear
attractive if the steam costs could be reduced.
The Phase II effort, as applied to the Cominco
Exorption process, was largely concerned with improving the ease of stripping
SO, from the absorbent. Although it appeared from the experimental results
that some reduction in steam costs might be realized, the overall process was
still not regarded as economical (see Part Three). As a result no Phase III
effort was required for the Cominco Exorption process.
(4) ' Mitsubishi Lime Process
The Mitsubishi Lime process afford a high quality
gypsum, suitable for sale as a constituent of wallboard, and for other less
widespread uses. Both capital and operating costs are low, and the latter appears
in Figure 37 as being almost independent of the sale of the product. However, if
the product could not be sold both a pollution problem and a large disposal cost
would result. The process is therefore highly dependent on producing a quality
product. This in turn requires the removal of nearly all of the fly ash contained
in the flue gas, and this must be accomplished upstream of the SO- scrubber.
A disadvantage of the process, although of minor
importance, is that of employing a slurry as the scrubber medium. All of the
other candidate processes have utilized solutions of salts for this purpose. A
more serious disadvantage involves the need for hauling the calcium oxide re-
quired for the process to the plant site. In the most unfavorable case this could
effectively nullify the credit obtained for the gypsum produced. It was concluded
that the process may well be economical only if the power plant to be serviced
were near a ready source of lime.
It was planned to pursue the Phase II and Phase III
investigation of this process if time was available.
-------
e. Sulfur Dioxide as a By-Product
i
Since gaseous SO, is an immediate product of three
of the candidate systems covered in this project, the cost of its conversion to
marketable forms of sulfur -containing compounds, viz. , liquid SO2, sulfuric
acid, and elemental sulfur, was of considerable interest. Conceivably, it may
be advantageous to consider any one of the three conversion routes for different
sets of circumstances. An order -of -magnitude cost analysis for each conversion
route is given in Appendix B.
D. SELECT PROCESS REFERENCES ;
This section contains brief statements about references which were
considered to be the most pertinent to the various processes evaluated. The
reference numbers correspond to the listing in the Bibliography, Part Six.
Fullham-Simon-Carves Process
111. A cost estimate is given.
202. Application of processes using ammonia for reducing the emission
of sulfur oxides from sulfuric acid plants is discussed.
237. A rather detailed account. Includes discussion on origin of process,
pilot plant at Fulham (2000 cu. ft. /min. ); 1940 half-boiler scale pilot
plant (60, 000 cu. ft. /min, ); laboratory work on the control of the
scrubbing process; and the North Wilford pilot plant. Results from
different pilot plants given. Process costs included.
358. General description of process given with some emphasis of process
446* Process chemistry is described. Applicability based on availability
of cheap ammonia liquor and the ammonium sulfate market assessed,
461. Paper gives a detailed account of the process. Gives information on
tht. various pilot plants, viz. , the Fulham, and the North Wilford
plants. Plant details and material of construction given for the
scrubbing system, filter press, autoclave and evaporator. Process
costs given.
566, A means for controlling the composition of the scrubbing liquid to
a narrow range of ratios of sulfite to bisulfite is described. Such
control is necessary to minimize the loss of ammonia vapor.
-------
621. Cost analysis for removing SO- from flue gases of a power plant
of 20,000-kw capacity considered, SO? content of gases treated
is 0.083 and 0. 30 percent. Estimates oased on 90% removal.
Process costs compared with those for other scrubbing methods.
639. Brief general description with flow chart. Numerous references
cited.
670. General description of process. Flow diagram and numerous
graphs giving various physico-chemical data, e.g., SO2 solu-
bility in water vapor pressures of SO2 and NH3 over ammonium
salt solutions, etc., are presented. Tables of pilot-plant data
given. Numerous references also provided.
673. Thorough coverage of process. Laboratory and pilot-plant
studies and operations discussed in detail. Extensive physico-
chemical data provided. Many references given.
682. General brief description of process. Flow diagram given.
Showa-Denko Ammomacal Process
390. Brief general description of process, including flow diagram.
Cominco Process
184. Special features, scope of application, and basic design principles
of the treatment methods used at Cominco1 s metallurgical opera-
tions at Trail are discussed. Some cost data are presented.
213. Application of the Cominco process as used at Trail, B.C., for
the treatment of gases containing relatively low concentrations
of SC>2, e. g. , as from coal- and oil-burning power plants, is
discussed. Process modifications needed to use the Cominco
process for such applications are discussed.
245. The development of the various sulfur-recovery processes at
Trail is discussed. Technical and economic considerations
are included.
255. Brief general description of process.
271. General brief description of process. Flow diagram given.
397. General discussion of the pollution problems created at Commco's
operations at Trail, B.C., and the control methods introduced to
solve such problems.
-------
445. Discussion of the application of the Cominco process for
treatment of the exit gae from a sulfuric acid plant at
Avonmouth, England. Scrubbing system is designed to
handle 32, 500 cu. ft. /min. Some details of the process
are given.
639. Brief general description of process. Flow chart given.
References cited.
654. Brief general description of process.
670. General description of process. Flow diagram given.
Operating data for various plants, e.g., lead-sintering,
zinc-roaster and acid plants presented. References pro-
vided.
672. Brief general description of process.
673. TVA studies, related to the Cominco process, which include
rather extensive pilot-plant work are discussed.
Cominco Exorption Process
24f>. General description of process and equipment.
670. Brief description of process.
Zinc Oxide Process
108. Some cost data for the process are given.
142. Process description and costs are presented.
255. Bri'?* general description of process. '
i
307. Brief general description of process.
35o. Br^ef generai description of process.
59*. Tire report, compiled by the people who developed the process,
is the most informative one available. A complete process de-
scription, including an account of aL. equipment components, is
given. Pilot-plant information and data provided. Estimates
on plant design, operation, and costs are discussed. Extensive
physico-chemical data related to process are given.
-------
621. Cost estimate of Zinc Oxide process given and compared with
similar estimates for other processes. A coal-burning power
plant of 120,000-kw capacity chosen as a basis. Coal con-
sumption of 1 Ib/kwh was used. Flue gas flow of 20,000,000
cu. ft. /hr with sulfur dioxide concentrations of 0. 083 and 0. 30%
were considered. A 90% removal of the SO, was assumed.
639. Brief discussion.
670. General brief description of process. Flow diagram given.
672. Brief general description of process.
673. Brief description of process.
Howden-I. C. I. (Cyclic Lime) and Mitsubishi
Simplified Lime Processes
15. General description of Mitsubishi Simplified Lime and other
processes.
111. The results of cost estimates for this process and the Battersea
process are presented. Study based on a coal-burning station
consuming 1 million tons per year. SO2 content of coal is 1-2%.
It is assumed that flue gas contains 90% of the sulfur and that SO
removal efficiency is 95%.
142. The results of a cost comparative study for the Howden-I. C. I.
(Cyclic Lime), Fulham-Simon-Carves, and Zinc Oxide proces-
ses are presented. A coal-burning power plant of 120,000-kw
capacity, using coal containing 1. 5 and 5 weight percent sulfur
was considered. The flue gas flow rate was taken as 20,000,000
SCFH with SO2 concentration of 0. 083 and 0. 30%.
237. Brief general description of process with emphasis on disadvan-
tages. Latter form basis for development of new process, the
Fulham-Simon-Carves process.
255. Brief general process description.
276. The development of the cyclic lime process is discussed. Be-
cause of the risk of scrubber blockages by calcium sulfate scale,
a study designed to ascertain the mechanism of scale formation
was conducted; the results are reported.
307. Brief general process description.
-------
312. Physical and chemical problems associated with the Howden-
I. C.I. process are discussed.
i '
328. The process is discussed in considerable detail. Description
and design principles of the complete process are presented.
Process chemistry, prevention of scaling, and operational
control are discussed. The pilot plant at Billingham is de-
scribed. The application of the process to large boiler plants
is also considered. A cost analysis for several cases, e. g. ,
for a power plant of 120,000-kw capacity, is also given.
356. General description of process with flow chart. Some dis-
cussion of historical development, particularly as related to
Batter sea process, is given.
358. A brief review of the process is given. Flow chart and some
cost data presented. Other processes, e. g. , Batter sea and
Fulham -Simon -Carves are also considered.
538. The occurrence of calcium sulfate crystallization and a method
of circumventing the problem are discussed.
621. Cost estimates of the Howden-I. C. I. process given and com-
pared with similar estimates for other processes. A coal-
burning power plant of 120, 000-kw capacity chosen as a basis.
Coal consumption of 1 Ib/kwh was used. Flue gas flow of
20, 000, 000 cu. ft. /hr with sulfur dioxide concentrations of
0. 083 and 0. 30% were considered. A 90% removal of the
SO 2 was assumed.
639. Brief general description with flow chart. Several original
references cited.
654. Bnel" discussion.
672. Brief general description.
Bartere«-»a Process
108. Liblii 'investment coyt of $11 to $15 pe-r installed kw, and
operating cost of 0. 3 to 0. o mill per kw-h.
151. Brier. Operating costs alone reported between $1.25 and
$1, 40 per ton of fuel, with totals probably double.
189. S ,ef. Shows scrubbing efficiency. Other systems discussed.
-------
192. Development data. Comprehensive description of power station
installation.
210. General discussion. Test data on catalytic oxidation of SO
using manganese and iron. 2
234. Brief description. Other processes discussed.
237. Brief discussion. Mentions disadvantages of requiring large
amount of Thames River water and contamination of the river
with CaSO. solution. Primarily Fulham-Simon-Carves paper.
255. Brief general description with flow diagram. Other processes
considered.
307. Brief description. Indicates operating cost of 12 to 15% of the
cost of delivered coal. Other processes considered.
356. General discussions concerning pilot-plant research, theoretical
considerations, the original and modified Battersea processes,
includes flow diagram. Also treats other processes.
Magnesium Hydroxide Processes
541. General description.
639. Brief description. Other processes discussed.
Magnesium Oxide Process
98. General discussion, equipment arrangements and design.
103. Equilibrium relations in the system MgO-SO^-H-O (acid region).
594. Shows oxidation data of SO2 and MgO calcination temperatures.
598. Provides pilot plant results of the equilibrium vapor pressure
of SO? over various magnesium bisulfite solutions, and the
application of the v'enturi gas scrubber.
639. Brief description. Other general discussions.
Manganese Oxide Process
5. SO2 used to recover manganese.
132. Equilibria in the systems
-------
346. Manganese ore slurry used for absorption.
424. Process description, some laboratory and pilot plant data.
439. SO, used to recover manganese.
Haenisch-Schroeder Process
183. General process description.
Wet Thiogen Process
465. General process description.
Ozone-Mn Ion and MnSO^ Processes
105. Laboratory data.
107. Reactions of SCX-O-, in solution of MnSO4>
209. Original laboratory work on the process.
210. Laboratory and pilot plant data. Negative results.
424. Excellent small pilot plant data.
448. Laboratory and pilot plant data.
639. Brief process description.
Sulfidlne Process
145. Describes pilot plant work. Unfavorable results. High
oxidation, high xylidine losses, excessive steam con-
sumption.
234. Process description. Operations on 5% SO, gas.
tt
366,, Plant description, flow diagram.
039. Brief process description. High xylidine losses.
-------
Basic Aluminum Sulfate Process
10. Describes development of process, flow diagram, plant
description.
41. Determination of the extent of the oxidation of the absorbed
sulfur dioxide by oxygen.
42. Laboratory packed column data.
224. Vapor pressure data. States that 1% SO- gas is lowest that
this process can economically treat.
639. Short description of process, unsuitable for gases containing
less than 1%
670. Brief description of process.
672. Short description.
Mitsubishi Amrnoniacal Liquor Process
15. General discussion and flow diagram. Other Mitsubishi
processes discussed.
Mitsubishi Manganese Oxyhydroxide Process
15. General discussion. Other Mitsubishi processes discussed.
Mitsubishi Lime Process
15. Brief description. Other Mitsubishi processes discussed.
Wisconsin Electric Power Process
344. General process description and design. Flow diagram.
Wisconsin Electric Power/Universal Oil Products Process
344. General process description and design, capital investment
and operating costs.
Combustion Engineering Process
341. Process discussion, development and large-scale operating
data, and cost estimates.
486. Process and apparatus description.
-------
Kanagawa Process
62. Brief statement about pilot plant unit in Japan.
257. Discusses salt water absorption system on board ship.
Wellman-Lord Process
429. General discussion with no information on proprietary special
treatment.
569. Belgian patent describing process.
Guggenheim Process
635. Brief description.
Diethylene Tnamine Process
634. Partial pressures of H_O and SO, over aqueous solutions of
diethylene triamine, triethylene fetramine, etc.
635. Laboratory study of absorption of SO? by aqueous solutions
of diethylene triamine and triethylene tetramine.
-------
PART THREE
LABORATORY EXPERIMENTATION RELATING
TO CANDIDATE PROCESSES - PHASE II
I. INTRODUCTION
At the completion of Phase I four processes had been selected for further
consideration in Phases II and III:
0 Zinc Oxide Process
• Cominco Exorption Process
• Ammonia-Hydrazine Exorption Process
• Mitsubishi Lime Process
II. LABORATORY EFFORT
The laboratory program related to each of these processes is discussed
below
A. THE ZINC OXIDE PROCESS
1. Johnstone Method
a. Process Description
In the Zinc Oxide process the flue gas is scrubbed
with an aqueous solution of sodium sulfite and sodium bisulfite. Zinc oxide is
mixed with the effluent liquor, forming insoluble zinc sulfite, and regenerating
soluble sodium sulfite which is returned to the scrubber. The zinc sulfite is
separated by filtration, dried, and calcined to produce zinc oxide, which is
returned to the process, and product sulfur dioxide.
Inasmuch as some oxidation occurs in the scrubber to
produce sulfate which cannot be readily calcined, the process includes pro-
visions for its removal. The effluent scrubber liquor is treated with insoluble
calcium sulfite, and the mixture is passed through a clarifier. The underflow
from the clarifier, which contains the calcium sulfite, is acidified with a
-------
portion of the product sulfur dioxide, theroby causing tho calcium aulfite to
dissolve. Calcium ion is thus made available for precipitation as calcium
sulfate, which is removed by filtration and discarded. The filtrate is treated
with lime to precipitate calcium sulfite, and it is then returned to the clarifier,
A flow diagram for the Zinc Oxide process is shown in Figure 10.
b. Process Reactions
Scrubber:
Na2S03 + H2O + S02 » ZNaHSOj (1)
Na2S°3 + 1/202 *" Na2S04 (2)
Liming Tank:
2NaHSO3 + CaO »• Na2SO3 + CaSOj ^ + H2O (3)
Gasifier:
CaSO3 + H2O + SO2 » Ca(HSO3)2 (4)
Ca(HS03>2 + Na2S04 » ZNaHSOj + CaSO4Jf (5)
Mixer:
2NaHS03 + ZnO -:~ 1 1/2H2O » Na2SO3 + ZnSO3-2 l/2H2ot (6)
Calcine r:
ZnSCy i. 1 /2H20 » ZnO + SO2 f + 2 1 /2H2O f (7)
c. Process Simplification and Improvement
(I) Introduction
Attempts to improve the Zinc Oxide process
included the following concepts:
« The use of monovalent ions other than
sodium ion in the scrubber as a means
of increasing scrubber capacity.
-------
• Lowering of the calcination temperature
required for the thermal decomposition
of zinc sulfite.
• Reduction of the extent of oxidation in the
scrubber.
Each of these categories will be discussed in turn.
(2) Scrubber Capacity
The capacity of the scrubber for SO, (see
£*
reaction 1) is dependent on the total Na concentration, C, expressed by Johnstone
as moles/100 moles H_O. Johnstone and Singh found that for C<3. 5 reaction
6 occurred exclusively when the rich scrubber liquor was subsequently treated
with zinc oxide. At higher values of C a second reaction, competitive with
reaction 6, also assumed importance.
6NaHSO3+ 3ZnO + H2O •» Na^Oy 3ZnSO3« 4H2O I + 2Na2SO3 (8)
Reaction 8 is undesirable because the sodium sulfite component of the double
salt does not release SO- under normal calcination conditions (£ 500 C).
The tendency of alkali metal ions to form
complex salts decreases with increasing charge density, so that Li for example,
appeared as a better candidate than Na for avoiding reaction 8. Accordingly an
experimental effort was planned in which Li , NH^ , CH-NH, and perhaps other
monovalent ions would be substituted for Na as a means of increasing scrubber
capacity. However, from data obtained in connection with the Phase III effort it
was subsequently shown that an increased scrubber capacity would probably not
improve the overall economics of the Zinc Oxide process. This arises from
the considerations that a certain minimum amount of liquid is required to
adequately wet the scrubber surfaces, and that for the quantity of influent gas
which must be treated in the practical case the capacity of the scrubber is
ample when C is 3. 5. Accordingly no experimental work was carried out in
this area.
-------
(3) Calcination
The temperature required to.release SO.
6
from zinc sulfite (see reaction 7) approaches 500 C during the final stages of
the calcination. Any substantial lowering of this temperature would be
expected to reduce fuel and equipment costs, and accordingly an approach
224
originally suggested by Johnstone et al. for reducing the cost of stripping
SO- from aqueous ammonia systems was investigated in connection with the
calcination problem. As applied to calcination the method involves the addition
of a solid acid, HA, to the calciner feed so that the acid anhydride SO- is
removed from a neutral salt (ZnA-) rather than a base (ZnO). The desired
reaction might be expected to occur at the melting point of the acid (compare
reaction 7):
ZnSO,- 2 1/2H,O + 2HA » ZnA, + SoJ+ 3 1/2H,O (9)
3 i, £•(.£'
If an acid were used in this manner, the counterpart of reaction 6, in which
the acid would be regenerated for further use, would also require demonstration:
2NaHSO3 + ZnA- + 2 1/2H-O • o» Na.SO.- + ZnSO.,' 2 1/2H2O? + 2HA| (10)
Ideally the acid should exhibit the following properties:
• Sufficient acid strength to promote reaction 9 at the melting point
of the acid
® No extraneous functionality, such as halogen, which might promote
side reactions
• A melting point in the approximate range 70° to 150°C. The lower
.limit implies that the acid should be solid at the scrubber temperature
(5G°C) to permit coprecipitation with zinc sulfite (reaction 10). The
upper limit is considered to be defined by the maximum temperature
at which steam might be conveniently utilized for the calcination step
(reaction 9).
*
The original concept of Johnstone is considered in II, B, relating to the Cominco
Exorption Process.
-------
• Non-volatility at the calcination temperature
• A relatively low molecular weight
In addition the acid should be inexpensive, and readily available in quantity.
The only acids which appeared to satisfy
the above requirements were &., u>-dicarboxylic acids containing four to eight
carbon atoms. The best candidate was considered to be azelaic acid (CgH./O,).
All of the acids of interest exhibit molecular weights of less than 200, which is
approximately the same as that of ZnSO,-2 1/2H9O (190). A preliminary
J L*
economic analysis indicated that if the weight of the calciner feed were increased
by a factor of two, and if the regeneration could be effected with steam at, say,
100 C, a substantial cost reduction could be realized. For a dicarboxylic acid
reaction 9 becomes
ZnSO3-2 1/2H2O H2A •» ZnA + SO2 f -I- 3 1/2H2Q (11)
Preliminary experiments with the candidate
acidb indicated that SO2 was partially evolved in accordance with reaction 11 at
the melting points of the acids. The presence of liquid water tended to promote
the reaction, perhaps through lonization of the acid. Inasmuch as the melting
points of all of the acids were above the normal boiling point of water, and since
little or no SO7 was evolved until the acids melted, it was considered that a
£•
mixture of candidate acids which melted below 100 C would offer the best
medium in which to retain the liquid water required to promote reaction 11.
Of various binary acid mixtures only those
involving azelaic and sebacic (G10H,gO4) acids melted in the desired tempera-
ture range. For these acids the eutectic occurred at approximately 80% azelaic
acid, and melted at 90° to 93°C. Subsequent experiments with the eutectic
mixture and water showed that SO2 was initially evolved at 90 , but that gas
evolution could not be maintained unless the added water was also volatilized.
From this observation, which is undesirable from an economic standpoint, and
the further observation that SO2 is readily liberated from aqueous K^SO^ by
-------
sulfuric acid at room temperature, it was concluded that an acid exhibiting
an ionization constant greater than that of azelaic acid (2. 5 x 10" ) would
better promote Reaction 11. However, no such acid was found which would
also meet the many requirements noted above.
Several quantitative experiments were
carried out with the candidate acids in which excess water, or acid, or both
were utilized. The best systems were those containing both excess water and
acid, but no more than half of the SO2 was evolved without boiling the water.
The following table summarizes the experiments conducted, where A = azelaic
acid, S = sebacic acid, AS = the eutectic composition, Z = ZnSOj* 2 1/2H2O,
W = water, and X indicates a four-fold excess of the constituent designated.
When no excess of acid was used the ratio of total acid to Z was always unity.
The evolved SO- is expressed as a percentage of that available in Z, and was
obtained after 20 minutes of heating.
System Bath Temperature (°C) %SO,
A - Z 115 - 125 27.6
A - Z - XA 115-125 25.2
A - Z - XA - XW* 95 - 100 35. 8
AS - Z 105 - 110 23.2
AS - Z - XW* 95 - 100 24. 3
AS - Z - XAS - XW* 95 - 100 50. 7
Based on water in ZnSO,« 2 1/2K-O
A clear melt was not obtained in any of the
experiments shown in the table, and it is believed that the insolubility of ZnA
may therefore be a limiting factor in the evolution of SO,, even if the water is
boiled. Although ZnSO3'2 1/2H2O may also be insoluble, this material is
subject to a heterogeneous reaction with the liquid acid (reaction 11), and
should therefore react completely unless product ZnA forms an insoluble
coating on the unreacted sulfite.
-------
Presumably an acid giving rise to a liquid ZnA would be desirable, even if
the latter were immiscible with the liquid acid.
Preliminary experiments involving various
ZnA derived from the candidate acids indicated that in all cases the zinc
compounds were infusible solids, which gradually darkened in color on heating
to 300°C. This may be attributable to the dibasic nature of the parent acid, in
that ZnA may be polymeric;
O
H|-0-C-(CH2)7— C-O-Zn-
O-C-(CH,)_-C-O-H
2'7
n
In the case of monocarboxylic acids, e. g. lauric acid, polymer formation cannot
occur, and the corresponding ZnA- are fusible:
zinc laurate, m. p. 128 C
However, from the standpoint of molecular weight, melting point, and other
properties, monocarboxylic acids are not suitable candidates for effecting
Reaction 9.
In summary, it is concluded from the above
results that the use of a weak acid no longer appears promising as a means of
lowering the decomposition temperature of zinc sulfite. If such an acid is used,
water will have to be boiled in order to efficiently remove SCX, and the reaction
may still not go to completion because of the insolubility of ZnA in the medium.
(4) Oxidation
594
According to Johnstone and Singh
at least 10 per cent of the available sulfite is oxidized to sulfate in the Zinc
Oxide process. The oxidation occurs mainly in the scrubber:
(2)
-------
The oxygen required for Reaction 2 is present in the flue gas, and the reaction
occurs most readily in dilute scrubber media, where the solubility of oxygen
is relatively high.
Reaction 2 is catalysed by transition metal
ions, particularly Mn4"1", Fe"1"1", and Fe+++. The number of d orbitals available
in these particular ions is such that the following type of structure can be
written, with SO," contributing electron pairs:
.Mn
-0 dl
It is postulated that Structure 1 is important to the oxidation process in that the
normally resonance -stabilized sulfite ion is rendered tetrahedral, so that the
electron pair shown on sulfur is readily available for oxidation:
1/202 + 3H20 — - •» Mn(H20)6 + SO4 (12)
As a result of Reaction 12, manganese ion is made available for further
eornplexmg:
(13)
Inua, Mn' functions as an oxidation catalyst.
5 94 644
It has been shown ' that phenols are
specific agents for the inhibition of the oxidation of SO-". It is believed that
this is to be attributed to the formation of a stable metal complex involving the
phenol, so that the metal ion is then no longer available for oxidation catalysis:
-------
H
-I- Mn(H2O)6
H20
Mn
OH.
-1 +
(14)
II
Structure II, which involves octahedral bonding (d^sp5) of the manganese, is
structurally related to the arene type of complex, of which dibenzene-chromium,
(C/H,)_Cr, is an example.
From the above discussion, it is concluded
that in principle it should be possible to effectively inhibit the oxidation observed
in the Zinc Oxide process through the use of phenols, and this in turn would
result in a substantial savings in chemicals cost (lime), waste disposal (CaSO,),
and equipment sizing. However, it was found in practice that the addition of
hydroqumone (a highly effective phenolic inhibitor) to the scrubber used in the
594
Zinc Oxide process had essentially no effect on the extent of oxidation.
A probable explanation is that rusted plant surfaces provided enough dissolved
iron so that the amount of inhibitor added was insufficient for complexing. In
this connection calculations have shown that the iron content of make-up water
is insignificant, and that the latter should therefore require a negligible amount
of inhibitor.
If the above explanation is correct, it should
be possible to inhibit the oxidation of sulfite through the employment of an
effective inhibitor used in conjunction with a scrubber which is constructed of
non-ferrous metals (or plastic) throughout. This argument presupposes,
however, that no fly ash is present. Inasmuch as fly ash normally contains
iron oxides as a major constituent, and since the scrubber medium for the
-------
Zinc Oxide process is maintained acidic, the amount of dissolved iron from
this source would require the use of a prohibitive amount of inhibitor, unleaa
the fly ash were nearly completely removed upstream of the scrubber.
The Wellman-Lord (Beckwell) process is
an example of an aqueous scrubbing process in which fly ash is largely removed
from the influent gas before scrubbing. The initial process step involves the
use of a water prescrubber, which removes the soluble portion of the fly ash,
most of the insoluble particulate matter, and any SO3 which had preformed as
a flue gas constituent. This method may also be applicable to the Zinc Oxide
process, and in the most ideal case no inhibitor would be required. It should be
pointed out, however, that if an inhibitor were required the process would have
to provide for a liquid effluent waste stream, so that soluble complexes
involving a cation of the type shown as Structure II could be removed from the
system. The present candidate Zinc Oxide process provides only for the
removal of a solid waste, in the form of calcium sulfate (see Reaction 5).
The use of a prescrubber for the Zinc Oxide
process was considered to be of sufficient interest to warrant an experimental
effort. Such an effort would involve a determination of the extent to which
oxidation occurs in a non-ferrous scrubber, when an aqueous prescrubber is
used to remove the soluble iron portion of the fly ash; and, the use of phenolic
inhibitors, if necessary, in the scrubber, as a means of combating oxidation
resulting from the presence of any iron which had not been removed by pre-
scrubbing. It was considered that the results of this investigation should be
generally applicable to aqueous scrubbing, inasmuch as many of the absorbents
for SO2 which have been used in aqueous medxa are oxidizable. Although some
fabrication of equipment was accomplished for conducting work in this area,
time and funding lor Phase II were not sufficient to work in the test program.
These studies were accomplished, however, in Phase IV {see Volume II).
2. Fluidized Bed Method
a. Process Description
From equations 1 to 7 it appears that the candidate
-------
and processing of SO.. In essence Zn acts as the most important of the
chemicals employed, since it is from ZnSO,' 2 1/2H-O that the SO, is finally
obtained. It was considered of interest, therefore, to determine whether ZnO
could be used for the direct absorption of SO-. If this could be accomplished
in a fixed bed, or preferably in a fluidized bed, neither Na nor liquid H,O
£»
would be required, and inasmuch as the only function of Ca is that of coping
with the oxidation of sulfite (Reaction 5), it was further considered that this
element might also not be required. This would be made possible by resorting
to the aqueous leaching of any ZnSO. formed from the insoluble sulfite.
Although water would be used in this step, the few bed volumes needed would
be small relative to the total water used in the Johnstone process, where in
the scrubber alone C has the relatively low value of 3. 5. The ZnSO. solution,
which could be of relatively small volume in view of the high solubility of
ZnSO., would be evaporated to dryness, and the solid returned for credit
against the ZnO makeup required.
The envisioned process was therefore one in
which the flue gas, after passage through a prescrubber to remove SO- and
fly ash, would be passed through a fluidized bed of ZnO for absorption of SO,*
The amount of water in the gas phase would be more than sufficient to permit
the formation of hydrated zinc sulfite. The latter would be calcined, as in
the Johnstone process, for the recovery of SO2 and the regeneration of ZnO,
which would be returned to the bed. Depending on the extent of oxidation, a
portion of the bed material would be leached with water to remove sulfate.
b. Process Reactions
Absorber:
ZnO + S02 + 2 1/2H20 - •> ZnSO3' 2 1/2H2O (15)
ZnS03-21/2H20 + 1/2O2 + 3 1/2H2O - *. ZnSO^ 6H2O (16)
Calciner:
ZnS03-21/2H20 - » ZnO + SO2 I + 2 1/2H2O T (7)
-------
c. Demonstration of Process Feasibility
(1) Introduction
2g
Previous data by the Bureau of Mines
have shown that solid zinc oxide is ineffective in absorbing SC^ from flue gas
at 130°C and at 330°C. However, it is postulated that the presence of an
adsorbed water layer on the zinc oxide surface is required in order to effect
the desired neutralization (Reaction 15), and that this condition was not fulfilled
at the relatively high temperatures employed in the referenced study. Water
is required not only as a highly polar medium in which to effect the reaction,
but also for the stabilization of the product through hydration.
(,2) Demonstration of Absorption
Initial studies were designed to show that
in the presence of sufficient water vapor, and at the prescrubber temperature
(50°C), appreciable absorption of SO- by fluidizcd zinc oxide can occur. No
attempt was made to approximate the SO- content (about 0. 3 vol %) of flue
gas at this time, but following the demonstration of absorption it was planned
to carry out additional experiments in the manner employed by the Bureau of
Mines.
The apparatus which was used for the
initial experiments is shown in Figure 38. In operation nitrogen gas and SO,
(The Uatheson Co., 99. 9%) were separately metered through calibrated
flowmeters, and the combined gas sparged through water which had been
previously saturated with SO^. In order to ensure that no liquid water entered
*he reactor the sparger was followed by a U-trap containing glass wool. As
a further precaution in this direction, the sparger liquid was permitted to cool
by the sparging of the gas, and was found to reach an equilibrium temperature
of about 21 C. Thus the gas entering the reactor was somewhat less than
saturated with water at all times. The U-trap was connected to the reactor
by means of tygon tubing, in order to provide the flexibility required for the
agitation of the reactor. The latter contained two course glass frits, with
-------
ro
SO,
f 18/9 Ball & Socket
Used Throughout
©Mercury Bubbler
(I) Flowmeter
(3) Water Sparger
® U-Tube Containing Glass Wool
® Flexible Tubing (Tygon )
©Reactor, 4 1/2" between Frits
(On center) x 1" 00
'inc Oxide Sample
Frit (Course •
Rubber Mallet
Thermometer
Alkaline Sparger
APPARATUS FOR ABSORPTION OF SULFUR DIOXIDE BY SOLID ZINC OXIDE
-------
the sample placed above the lower frit. Agitation was required to prevent
channeling, and was provided by means of a rubber mallet, which in turn
was attached to a Burrel wrist-action shaker. The reactor was wrapped
with electrical heating tape for runs conducted above ambient temperature.
The exit gas from the reactor was passed through aqueous sodium hydroxide
to absorb unreacted SO,.
b
Three different grades of zinc oxide were
examined. Two of these were obtained from the New Jersey Zinc Company,
and were designated as Horse Head Kadox —15, 99. 7% pure, 0. 11 micron
mean particle size, and Horse Head XX-504, 99. 6%, 1. 5 microns. The other
material was zinc oxide powder, reagent grade, 99. 0%, of unspecified particle
size, and was obtained from the Allied Chemical Company. From a compari-
son of this material with the other two under fluidized bed conditions it appeared
that it was probably intermediate in particle size between the New Jersey Zinc
*
Company products.
The experimental data for a complete run
for K_adox —15 at room temperature are shown in Table 68 and a plot of con-
version to ZnSCy 2 1/2H2O vs time is shown in Figure 39. It will be observed
that essentially complete absorption was effected under the conditions employed,
Very slight balling of the material was observed to occur during the initial phase
of the run, and after 75 minutes the sample was removed from the reactor and
*
Although Kadox —15, of 0. 11 micron mean particle size, was used in most of
work to be described, it is considered doubtful that such finely divided material
is required. Johnstone and Singh594 observed that their freshly regenerated
oxide, which was approximately 30 to 140 mesh, was highly reactive in the
sense that complete dissolution of the oxide occurred in aqueous sodium
bisulfite solution (see Reaction 6) within "a few seconds. " On the other hand,
they noted that a commercially obtained oxide required at least two minutes.
In the present work it was noted that all of the samples of zinc oxide described
above required about three minutes for essentially complete solution in aqueous
sodium bisulfite. It would appear, therefore, that commercial zinc oxides may
form a monomolecular coating of the less basic zinc carbonate, or perhaps the
peroxide, on long exposure to air, and that in consequence freshly regenerated
oxide of relatively large particle size may be more reactive than any of the
oxides considered here.
-------
TABLE 68
CONVERSION OF SOLID ZnO (KADOX-15) TO ZnSOj-2-1/2
AT ROOM TEMPERATURE - EXPERIMENTAL DATA
Zinc Oxide: 0. 856 g
Time
(mm)
-30
-15
0
15
25
35
55
75
85
105
130
155
185
215
245
275
305
330
™
SO Absorbed
2 (8)
_
0.000
0.010 (H20)
0. 183a
0.261
0. 339
0.437
0.509b
0.570
0.654
0. 753
0.825
0. 922
0.979
1.025
1.051
1. 071
1.077
1. 154
Conversion to
ZnSO,-2-l/2H,O (%) Remarks
J £,
Start
N2
N2 + H2O
15.8 N, + SO, + H.O
c* £* £»
22. 6
29.4
37. 9
44.0
49.4
5(3.6
65.3
71.5
79.9
84.8
88.8
91.2
92.8
93.4
100.0 Theoretical
^reabsorbed H2O (0.010) not included.
Sample removed and ground in mortar, subsequent values
corrected for mechanical loss incurred.
-------
100
95
90
85
80
75
70
65
* *°
55
50
45
40
35
30
25
201
15
10
5
ett
is
1
J
0 30 60 90 120 150 180 210 240 270 300 330 360 390 420
Time (Minutes)
CONVERSION OF SOLID ZnO < KADOX -15) TO ZnSOj^ 1/2
AT ROOM TEMPERATURE
Figure 39
-------
ground in a mortar. No additional grinding was considered necessary, and
the fact that Reaction 16 proceeded essentially to completion, without the
necessity for repeated grinding was interpreted as indicating that the zinc
sulfitc coating is somewhat porous. As indicated in Table 68, the run proper
was preceded by a fifteen minute period involving sweeping with nitrogen,
followed by an additional fifteen minute period with nitrogen which had passed
through a water sparger. This pretreatment was designed to provide a surface
layer of water on the zinc oxide before the introduction of SO~. At the time of
introduction of the SCX, the water sparger was replaced by a sparger containing
water which had been saturated with SO?. The gas flow rates and gas compo-
sitions which were used are shown in Table 69. The water content was de-
termined by weighing the sparger initially and after sparging. The final product
from the initial run was a free-flowing powder, essentially identical in ap-
pearance with the starting material.
Additional experiments were conducted in
which the effect of water vapor, particle size, and temperature on the absorption
were investigated. It was found that in the absence of water vapor little or no
absorption occurred, and that the absorption decreased with increasing tempera-
ture at constant humidity. The absorption rate also decreased with increasing
particle size of the absorbent, indicating that reaction occurred primarily at
the particle surface.
The experimental results described above
cannot be directly related to the absorption characteristics of zinc oxide for
SO_ contained in flue gas. Additional experiments were carried out, therefore,
in which a fluidized bed of zinc oxide was used under the experimental conditions
utilized by the Bureau of Mines in their screening program for metal oxides,
but with the following modifications:
• The zinc oxide study was conducted at S 50 C,
instead of 130°C and/or 330 C.
• The zinc oxide particle size was smaller than
that of the oxides prepared by the Bureau of
Mines.
-------
TABLE 69
GASEOUS FLOW RATES AND COMPOSITIONS USED
IN REACTOR SHOWN IN FIGURE 38
N2
N2
H20
N2
Hz0
SO,
N2
S02
Flow (25°C and 752 mm)
1/hr moles /hr g/hr
37.7 1.680 47.1
37.7 1.680 47.1
1.00 0.044 0.80
30.5 1.360 38.2
1.01 0.045 0.81
7.20 0.321 20.5
30.5 1.360 38.2
7.20 0.321 20.5
Composition
Wt-%
100
i
98.4
1.6
64.2
1.3
34.4
i
65.1
34.9
Vol-%
100
97.3
2.7
78.8
2.6
18.6
80.9
19.1
-------
• Fc • simplicity al tlud time tho ga& utilized
in the present i>tudy was 0.27 vol-% SO- in
N . >n the Bureau of Mines study CO., and
>. (but no NO ) was also incorporated in the
gas. x
• The space velocity fvol of gas (S. T. P.) per hr
per vol of absorbentj of 1050 used by the
Bureau of Mines corresponded to that employed
in the present study for the dry gas (SO- in N,).
Following the water sparging operation the space
velocity was somewhat higher.
The apparatus which was used is shown in
Figure 40. The dry gas was metered at a space velocity of 1050 hr" into the
water sparger, which was maintained at approximately the selected absorption
temperature by means of a water bath. Glass wool was placed directly in the
sparger unit to prevent the introduction of liquid into the reactor. The upper
part of the sparger was wrapped with heating tape, which was extended to include
the entire reactor. The latter was similar in design to that shown in Figure 38,
but was made longer* so that gas entramment would result in a minimal deposition
of solids at the upper frit. The reactor was agitated by means of a vibrating
table. A thermometer well was fabricated so that the bulb of the inserted
thermometer was in the zone just above the surface of the zinc oxide charge, and
•vas therefore continuously bathed with fluid zinc oxide during the run. Metallic
mercury was used for heat transfer in the thermometer well. Unreacted SO-
in the exiting gas was absorbed in standard 0. 1 N I_, and the excess iodine
titrated with standard 0. 1 N Na-S-O-. The iodine trap was analyzed at selected
times during the run.
Preliminary experiments were conducted
to determine the efficiency of the iodine trap under operating conditions. In
the absence of excess potassium iodide, 25. 5% of the contained iodine was lost
by volatilization when nitrogen was sparged through the solution for one hour at
a space velocity of 1050 hr" . This loss was decreased to 1.66% through the
addition of 5 g of potassium iodide for each 20 ml of 0. 1 N I2 employed, and to
0. 54% when a 10 g excess was used. The latter amount was always used during
the run, with a correction included for the small iodine loss incurred. Another
-------
Note:
Upper portion of water sparger and entire
reactor wrapped with electrical heating tape
§ 28/15 Ball &
Socket
00
® Mercury Bubbler
Flowmeter
Flexible Tubing flygon)
ter Sparger
lass Wool
I Water Bath
Heater
(D Thermometer
® Reactor. 15" Between Frits
(On center) x 1" 00
Zinc Oxide Bed
Glass Frit
Vibrating Table
Thermometer Well and
Thermometer
Iodine Sparger
FLUIDIZED BED REACTOR SYSTEM
-------
experiment was designed to show that all of the available SO- was absorbed
when 0.27 vol-% SO^ in N, was spargi d through the iodine solution at a space
velocity of 1050 hr for one hour. f h.it no SO, was lost was shown by back
titration with tluosulfatc, which indicated a. value of 0.27 vol-% SO- in the gas,
in agreement with the known SO, content of the gas as previously determined
b
by mass spectral analysis.
A rather complex run which was carried
out with fluidized Kadox —15 is shown in Figure 41, and the corresponding
experimental data are given in Table 70. Gaseous flow rates, compositions
and wet space velocities used during the run are shown in Table 71. The water
content of the gas as indicated in the table is less than that expected for
saturation at either 35 or 50°C. This is to be attributed to the fairly rap;.d
rate of sparging, which would both lower the temperature of the water to
some extent (compare the above section) and would provide insufficient contact
time for saturation. It may be noted that the water content of a typical flue
gas without water prescrubbing (which would tend to increase the contained
water) is approximately 7. 25 vol-%. Inasmuch as saturation of the gas with
water was not achieved in the experiment to be described, the experimental
data relating to the extent of SO, absorption may be considered as conservative
£•
in relation to those which would be expected in practice. The run proper, which
was initiated at 35 C, was preceded by a water pretreatment period of 105
minutes during which nitrogen gas, which had been sparged through water at
35°C, was passed through the bed. No mechanical difficulties were experienced
with the bed during this period or subsequently, and the bed remained highly
fluid at all times. Channeling was effectively prevented through the use of the
vibrating table.
The first fifteen minute period of the run
was allowed for the establishment of equilibrium conditions, including saturation
of the water sparger with SO-. The iodine trap was incorporated into the system
at the end of this period, and was removed one hour later for back titration of
the excess iodine with sodium thiosulfate. Throughout the run the iodine trap
was similarly incorporated into the system for one hour periods, and at the
-------
§
'•i.
g
ICQ
95
90
85
80
75
70
£5
60
55
50
35°C-»— 1—-50 °C
1 \
-
-
A
\
\
\X<
\
D
\/
y
\
\
/
r
\
\
\
C\
\
<
r^
F
\
\
_
v
^
\
>
ris
i
\
\
H \
t
\
,s
•
G 50 100 150 200 250 300 35C 400 450 500 550 600 650 700 750 800 850 900 950 1COO
Time (Minutes)
CONVERSION OF SOLID ZnO ( KADOX-15) TO ZnSOj-21/2H20 IN A FLUIOIZED BED REACTOR
-------
TABLE 70
REACTION OF FLUIDIZED ZnO (KADOX-15) WITH 0.27 VOL-% SO
IN N2 AT 35°C AND AT 50°C; EXPERIMENTAL DATA
Zinc Oxide:
15. 8 g
Space velocity of dry gas: 1050 hr
-1
J. illiC
(mm)
-105
0
15
75
185
300
380
460
570
630
690
775
Reactor Water Bath
35
36
35
35-1/2
35-1/2
35
35
50
50
50
49
49
35
33
34
34
34
33
33
48
49
48
48
48
t?v^r_ rvuaui ucu
2 (%)
_
-
-
100.0
100.0
100.0
-
100.0
83.9
93.5
98.0
77.9
Remarks
Start
N2 + H2C
N + SO
it b
Shutdown
of bed.
Shutdown
resumed
Shutdown
Shutdown
Shutdown
>
+ H20
for inspection
overnight;
at 50°C.
for 3 hours
overnight
and maintaine<
*"\
835
895
49
49
49
49
78. 8
67. 3
system at 50°C for one
hour.
Shutdown and maintained
system at room temp.
for 50 minutes.
Total pickup (SO2 + H2O):
6. 6 g.
-------
TABLE 71
GASEOUS FLOW RATES. COMPOSITIONS AND SPACE VELOCITIES
USED IN REACTOR SHOWN IN FIGURE 40
Flow (S. T. P.) • Composition Space Velocity
{1 /hr) (Vol-%)
N2
N2
aH2O sparger at 35°C
bH2O sparger at 50°C
34.1
1.43
0.09
95.7
4.01
0.25
)
)
-
1095
34.1
2.14
0.09
93.9 )
\
5. 90 )
0.25 )
1118
-------
times indicated by the data in Table 70.
\nalyscs at 75, 185, and 300 minutes
showed that complete absorption of the SO, was occurring. The bed was
£
highly fluid during this time, and inasmuch as it was not possible to de-
termine by inspection whether the bed contained liquid water, the run was
momentarily interrupted after 300 minutes. It was found, however, that no
liquid was present. This was confirmed by noting that the weight gain of the
reactor (2. 3 g) corresponded almost exactly with that expected (2. 2 g) for
the conversion of the total £O_ absorbed during the first 300 minutes to
ZnSO3-2-l/2H2O.
At 380 minutes (shown as A in Figure 41),
the run was shut down overnight, and because complete absorption of the SO-
had occurred up to this point it was decided to resume the run at 50 C. An
analysis after 460 minutes showed that complete absorption was still taking
place. However, at 570 minutes (B) absorption was incomplete, and from
the results obtained by the Bureau of Mines in their screening program, it
was considered that the run could be considered as having been nearly
completed at this point. In the Bureau of Mines study a given run was continued
until the candidate material no longer absorbed at least 90% of the incoming SO-.
In general it was found that absorption was complete until the absorbent was
spent, at which time the absorption dropped sharply. In the present study
it appeared necessary to obtain only one additional point C, beyond B, so
that a line BC could be used to establish the point D, at which a minimum
of 90% absorption no longer occurred. From these data one could then calculate
the total absorption, in gram per 100 grams of absorbent, for comparison with
similar data for the various oxides which were screened by the Bureau of Mines.
Surprisingly the next two points, at E and
F, showed appreciable increases in absorption. These increases were well
beyond experimental error, and a review of the data indicated that the only
unusual circumstance attending points B and E was that in each case the run
had been interrupted; at B for 3 hours, and at E overnight. No interruption of
-------
the run occurred at F, and the next point, G, showed & considerable d.ecreaio
in absorption.
Two hypotheses were considered to explain
the observed phenomenon. One of these would involve the rupture, on tempera-
ture cycling between 50°C and room temperature, of the zinc sulfite coating
which had formed on the surface of the zinc oxide during the run, thereby
exposing a fresh surface of the oxide to further reaction. It was noted in
this connection that the interruption of the run at A ha;d resulted in nearly
complete cooling of the solid after only 15 minutes. If this explanation were
valid, rapid cooling of the solid at G should show a pronounced effect on the,
absorption curve, and the otherwise expected point H would not be realized.
The other hypothesis is based on the fact that zinc sulfite, unlike the oxide,
is slightly soluble in water (0.16 g/100 g at room temperature)* This implies
that the water present on the solid surface is saturated with zinc sulfite, and
that a digestion process might occur in which the zinc sulfite coating continuously
dissolves and reprecipitates as a separate crystalline phase. The exposed zinc
oxide surface would then become available for furthet absorption of SO,. The
digestion process would be favored by an increase in temperature. If this
explanation were valid, the rapid cooling of the solid would probably not result
in subsequent enhanced absorption, since the digestion process would be
expected to be quenched in this case. The enhancement of absorption would
rather be promoted by interrupting the run, but at the same time maintaining
the zinc oxide bed at 50°C until the run was resumed.
The second of the hypotheses considered
above was examined first. At point G the run was interrupted, and the system
was maintained at 50 C for one hour before the run was resumed. The position
of point J, which was subsequently determined, was interpreted as indicating
that the digestion process is at least an important factor.
The importance of the rupture of the zinc
sulfite coating was examined by quickly cooling the bed at point J through the
use of cold air, and subsequently permitting the system to stand for one hour
-------
before resuming the run. Heating the - ystcm to 50°C was accomplished
during the final ten minutes of this oru hour period. The next experimental
point, K. , indicated that temperatu.c cycling was ineffective in promoting
the subsequent absorption of SO-, and therefore that the rupture of the zinc
aulfite coating is probably not an important factor.
A few final remarks may be of interest
in connection with the experiment described above. The observation that the
slope of BE > slope of EF > slope of GJ may perhaps be attributed to the
longer time available for digestion at B (3 hours) and at E (16 hours) even
though the system was at room temperature during most of the time following
B and E. Although the system was also interrupted after 300 minutes, no
absorption effect was evident in this case, since complete absorption was
still being accomplished at this time. Finally, it may be noted that point D,
as indicated in Figure 41,is probably realistic, inasmuch as the slope of the
line BC is approximately the same as that of the experimentally determined
lines FG and JK.
The theoretical SO- pickup per 100 g of
absorbent per hour for the system discussed above is 1.66 g. As noted
above, this value was precisely realized at the 300 minute point. At point
D the SO2 pickup should have been 515/60 x 1.66 + (550-515)760 x 0. 95 x 1.66 =
15.2 g. This number may be considered as minimal, inasmuch as the digestion
process noted above would normally be expected to enhance the pickup in the
practical case, where only part of the spent oxide would be removed for
regeneration and the rest would be recycled. To the extent that time would
be required for the recycle process, digestion would be expected to occur,
so that the recycle feed would possess greater activity than the fresh effluent
from the absorber. Table 72 gives a comparison between the results of this
study and that conducted by the Bureau of Mines for sodium aiuminate'and for
alkalized alumina. The comparison is necessarily crude, because of the
variables involved (temperature, particle size, space velocity, humidity, etc.),
but does serve to indicate that zinc oxide should be considered as a candidate for
the absorption of SO- in a fluidized bed application.
-------
TABLE 72
SULFUR DIOXIDE ABSORPTION BY SELECTED SOLID ABSORBENTS
a
Absorbent
NaAlO2
Al 0 )
2 3 }
Na20 )
ZnO
NaAlO,
£»
A12°3 1
Na20 )
ZnO
ZnOb
Purity
(wt-%)
96
73
25
-
96
73
25
-
99.7
Bulk Density
(g/cc)
0.90
0.54
-
0.90
0.54
-
0.48
SO2 Absorbed
(g/IOO g absorbent)
18
10
17
Temp
-J!£L
130
330
50
Particle size and space velocity: this study, 0. 11 microns and 1118 hr ,
respectively; Bureau of Mines, 8-24 mesh and 1050 hr***.
This study; all other data taken from the Bureau of Mines study.
-------
The total pickup of water and SO2> 6. 6 g, at the end
of the experiment (see Table 70), corresponds to a conversion of 31. 0% of
the zinc oxide to ZnSC>3' 2-1/2 H2O. The final solid was a free flowing
powder, which did not appear to have adsorbed appreciable surface water. A
sample of the material was found to lose 9. 3% by weight when heated to 105 C
for 70 minutes. Commercial ZnSO.-2-l/Z H^O lost 13.8% under the same
conditions. In each case some water of hydration was probably lost, in addi-
tion to surface water.
Some of the more important conclusions to be derived
from this initial Phase II study are the following (see Volume II of this report
for additional results associated with the development of this concept and for
additional conclusions based on these later results. )
• A fluidized bed of zinc oxide is effective at 50 C for the
absorption of SO_ from a carrier gas containing 0. 3
vol-% of SO2. *•
• The absorption process is favored by small particle size,
low temperature, and the presence of water vapor. If the
latter is not present, no absorption occurs, on the other
hand, a liquid water phase is not required for absorption.
• Inasmuch as the water content of the gas utilized in the last
experiment considered above was less than that required for
saturation of the gas, the absorption data obtained from this
experiment are considered as conservative in relation to
those which would be expected in practice.
• The zinc sulfite coating formed during absorption may be
porous to the further penetration of SC^. An alternative
explanation involves a digestion process in which the zinc
sulfite coating slowly dissolves in surface water and
reprecipitates as a separate crystalline phase. Both
concepts may be valid.
• A rough evaluation of the absorption characteristics of zinc
oxide indicates that it compares favorably with sodium
aluminate and with alkalized alumina for the absorption of
so2.
-------
• The temperature and humidity of a flue gas following an
aqueous prescribing step is precisely that required
for effecting the absorption of SO2 by a fluidized bed
of zinc oxide; i. e., 50°C and a water content which is
near the saturation value.
(3) Calcination
The calcination step for the zinc oxide
bed method is identical with that used in the Johnstone method, as discussed abovi
(4) Oxidation
The method envisioned for the separation of
zinc sulfite from its oxidation product, zinc sulfate, involves leaching of the
sulfate with water, as noted above. Among the relatively inexpensive absorbents
for SO2, zinc appears to be a particularly good candidate for the separation of
the sulfite and sulfate in this manner. In the case of alkali metals, for example,
both the sulfites and sulfates are soluble in water, and this is also true to a
lesser degree of magnesium. On the other hand both calcium sulfite and calcium
sulfate are insoluble.
The absorption step for the fluidized bed
process is accomplished in the absence of liquid water, and it is of interest to
consider whether appreciable oxidation is to be expected under these conditions.
In aqueous systems the oxidation process is frequently homogeneous, and
typically involves the reaction of dissolved oxygen with the diasolved absorbent
(e. g., Na_SQ~ in the Johnstone process) to give a soluble product (e. g., Na-SO^),
Of special importance is the fact that the catalysts (e. g. , Fe ) which promote
the lew temperature (50°C) oxidation process are also water-soluble. In the cast
of the fluidized bed any catalyst present will necessarily be in the form of a solid.
and no leaching of iron from fly ash not retained by the pr esc rubber will occur.
It is also questionable whether appreciable oxygen will dissolve in an adsorbed
monolayer of water on the zinc oxide surface. It is believed, therefore, that
inasmuch as the conditions leading to homogeneous oxidation catalysis in
aqueous scrubbers will not be present under fluidized bed conditions considerably
-------
leas oxidation will occur in gas -solid systems.
3. Attempted Synthesis of Zinc Pyrosulfite
Some experimental work was carried out on the program
in an attempt to prepare zinc pyrosulfite, ZnS-O-'xH-O. A compound of this
composition has apparently not been reported. Zinc pyrosulfite would offer the
advantage over the simple sulfite of affording two moles of SO, on calcination:
ZnS205- x H20 - •» ZnO + ZSO + x H2O (17)
ZnSO' 2 1
3 /2H2O . «• ZnO + SO2 + 2 1 /2H2O (7)
The existence of pyrosulfite ions in aqueous solutions of potassium bisulfite
is well established, and forms the basis for the Wellman-Lord process for
the removal of SO- from flue gas. In this process pyrosulfite separates from
a cool aqueous solution as a crystalline solid:
2KHSO3— *• K2S205+H2O (18)
Zinc oxide is a relatively weak base, and it is therefore
of interest to consider whether the neutralization product zinc sulfite might be
expected to combine with an additional mole of SO- to form the pyrosulfite.
In general this type of reaction occurs only with salts derived from strong
bases, for example potassium sulfite combines with additional SO- to form the
bisulfite in solution, or the pyrosulfite as a water-free solid (Reaction 18). In
the case of ammonia, which IB comparable in base strength with zinc oxide,
the bisulfite forms in aqueous solution, but attempts to isolate either this
compound or the pyrosulfite by solvent removal result in a loss of SO2. This
would suggest that either zinc pyrosulfite might be non-isolable or, if isolated,
would exhibit a high decomposition pressure of SO2 at relatively low tempera-
tures. The possibility of preparing the compound was nevertheless considered
of sufficient practical interest to warrant a preliminary investigation.
In initial experiments designed to prepare zinc pyrosulfite
-------
aqueous solutions of zinc chloride and potassium pyrosulfite (which exists in
solution as the bisulfite) were mixed at room temperature, and the resulting
solutions cooled to 10°C to induce crystallization: i
ZnCl2 + 2KHS03+(x-l)H2O — •» ZnS^g-xH^O I + 2KC1 (19)
However no precipitate formed, even from solutions nearly saturated with the
starting materials at 10°C. It was concluded, therefore, that either the
desired compound does not form, or it is highly soluble in aqueous media.
That solid zinc pyrosulfite does not form from solid
zinc sulfite and gaseous SO- in the presence of water vapor was shown in
6
connection with work reported above relating to the absorption of SOg by
fluidized zinc oxide. From Figure 39 it appears "that solid zinc oxide absorbed
approximately 94% of the SO2 required for the formation of ZnSO3= Z 1/2 H20,
with no tendency exhibited toward the absorption of a second mole of
Additional experiments were conducted in which an
attempt was made to form zinc pyrosulfite in the following manner (compare
Reactions b and 18):
4KHSO3 + ZnO + (x-2) I^O - s» ZnS^g- x H2
-------
TMs was observed at C=6, S«*5, S/C=0.83, where S is the concentration of
available SO2, in moles/ 100 moles H^O, and at C =2, 5=1.67, S/C=0.83.
Although the observed precipitate may have contained pyrosulfite, it was
considered that the occlusion of zinc oxide would render this system value-
less in the practical case.
From the results of the above experiments it was
concluded that zinc oxide may well be too weakly basic to form an isolable
pyrosulfite. Accordingly, work in this area was terminated.
B. THE COMINCO EXORPTION PROCESS
1. Process Description
The Cominco Exorption process involves the scrubbing
of flue gas with an aqueous solution of ammonium salts. The off-stream liquor
is heated for the liberation of SO2> the spent liquor then being returned to the
scrubber. To the extent that oxidation occurs in the scrubber ammonium
sulfate is also produced. The process was used early in World War II when
ammonia was in short supply, so that other ammonia-based processes which
converted all of the SO- to ammonium sulfate (e. g. , the Fulham -Simon -Carves
and Cominco processes) were less attractive. A flow diagram of the Cominco
Exorption process is shown in Figure 8.
2 Process Reactions
Scrubber:
2NH4HS03 (21)
(NH4)2S03 + 1/202 » (NH4)2S04 (22)
Heater:
2NH4HS03 •» (NH4)2S03 + SO2 f + H2of (23)
-------
3. Process Simplification and Improvement
a. Introduction
The Cominco Exorption process involves somewhat
higher capital and operating costs than those for the other candidate processes
(see Figures 36 and 37 ) and this can be attributed in large part to steam and
equipment costs associated with the desorption of SO,. Accordingly, emphasis
in the laboratory was placed on effecting a marked reduction in steam re-
quirements through the use of an acidic type of additive, as considered above
for the Zinc Oxide process. However, the physical properties required of an
additive for the Cominco Exorption process differ markedly from those required
in the case of the Zinc Oxide process. In order to avoid the formation of a slurr
in the scrubber, an additive for the Cominco Exorption process should melt beta
the scrubber temperature of 50°C, and in order to maintain capacity of the
ammonia-based scrubber liquor it should preferably be nearly insoluble in the
scrubber liquor at this temperature. However, the additive should be highly
soluble (but non-volatile) at the stripping temperature, inasmuch as the
promotion of SO^ -desorption at approximately 90° to 100°C occurs through
solution of the additive. The following equations illustrate the method:
NH4HS03 + HA 90 to
50
NH4A + S02 + H20 "v ^» NH4HS03 + HA» (25)
Another important criterion for a candidate acid additive
is that the following relationship should be approached as closely as possible:
K K = 10~4*4
a s
where: K is the ionization constant of the acid
and K is the molal concentration of the un-ionized
portion of the acid in solution.
The above equation was derived by Johnstone et al.224 for a
-------
partially soluble or partially miscible ?cid of the type under discussion. When
the relationship holds the optimum buffering action in the scrubber commensu-
rate with maximum SO, absorption >a realized.
b. Desorption of Sulfur Dioxide
The feasibility of using an acid as an additive for
the purpose of lowering the steam requirements for the Cominco Exorption
process has been demonstrated by Johnstone et al. who measured
the partial pressure of SO2 in ammonium sulfite-bisulfite systems both in the
presence and absence of selected acids. A significant increase in the vapor
pressure of SO2 was observed at 90°C with solutions containing, for example,
valeric or caproic acid. However, these systems were not considered practical
because of various undesirable properties (volatility, high melting point, etc.)
of the acids.
As a result of a survey of the various acids which
are available in bulk at low cost (OP&D Reporter), it appeared that in all
probability no single acid would possess all of the properties considered above.
In particular, high boiling acids which exhibit little or no volatility at the
temperature of regeneration of SO, invariably melt higher than 50°C. It was
believed, however, that a mixture of acids might exhibit a eutectic melting
well below this temperature, so that over a range of acid compositions the
overall scrubber medium would appear as a two-phase liquid system.
Two acids were selected as initial candidates
for effecting Reaction24: azelaic acid, HOCO(CH2)7COOH, because of its
ready availability, low price (36 cents/lb), non-volatility (vapor pressure of
1 mm of Hg at 178°C), and low water-solubility (0.24 g/100 g H2O at 20°C,and
2.2 g/100 g KLO at 65°C); and HET (hexachloroendomethylenetetrahydrophthalic)
acid, because of its relatively low melting point (70°C), low water-solubility
(0. 35 g/100 gH2O at 23°C) and high miscibility with water at 96° to 97°C.
It was considered that if either of these acids appeared promising additional
effort would be justified involving the use of relatively low melting acid
mixtures.
-------
The experimental results obtained with azclaic
and HET acids indicated that the major problem area was precisely that
previously encountered in connection with the use of acid additives as an aid
to the calcination of zinc sulfite (see above); namely, that the desorption of
appreciable SO- from the rich solution (Reaction 24) required that a relatively
large quantity of water be volatilized. Although some reduction in steam
requirements appeared to be feasible, the extent of reduction did not appear
to justify the use of the acid.
In subsequent work with oxalic acid, (COOH),,
it was found, as had been anticipated, that the high acid strength of this acid
permitted the relatively efficient desorption of SO2. However, the solubility
of the acid was such that the Johnstone relation considered above did not hold,
and as a result the desorbed solution showed little capacity for the further
absorption of SO,.
£»
C. THE AMMONIA-HYDRAZINE EXORPTION PROCESS
i. Introduction
The Ammonia-Hydrazine Exorption process was con-
ceived at Aerojet, and is similar in principle to the Cominco Exorption process,
but potentially free of the most serious difficulty experienced with the latter;
namely, the high steam requirements for the desorption of SO,. Because of
It
the much greater water solubility of hydrazine salts relative to ammonium
salts, it was considered that the scrubber liquor for the hydrazine system
would contain relatively little water. In fact, the melting points of hydrazine
ajt
salts in general are sufficiently low that in the limiting case the scrubber
liquor might be comprised of a low-melting eutectic or near-eutectic mixture
of salts, with water present only to the extent of an equilibrium amount arising
from the flue gas. Under these conditions steam requirements for the desorption
step would be minimal, with very little water being volatilized during desorption.
* - - ,o
For example, hydrazine sulfate melts at 87 C; compare ammonium sulfate,
which melts with decomposition at 513°C.
-------
The process also appeared attractive that anhydrous hydrazine, a process
intermediate, would represent a sale-iblo product to the extent that a market
exists or could be developed, with . i*plus hydrazine being returned to the
process. Makeup hydrazine would be in the form of a dilute aqueous solution,
which is commercially available, and which is considerably less expensive
than the anhydrous product. Because of the small quantity of makeup likely
to be required, the amount of water introduced into the scrubber in this manner
would not appreciably affect the overall water content of the scrubber liquor.
2. Process Description
The Ammonia -Hydrazine Exorption process involves
the scrubbing of flue gas with a concentrated solution of hydrazine salts. The
off-stream liquor is heated for the liberation of SO2, and the spent liquor is
then returned to the scrubber. To the extent that oxidation occurs a portion
of the spent liquor is ammonolyzed in liquid ammonia at room temperature to
produce ammonium sulfate and anhydrous hydrazine. A flow diagram of the
Ammonia -Hydrazine Exorption process is shown in Figure 31.
3. Process Reactions
Scrubber:
i
(26)
Heater:
Ammoniator:
+ 2NH, » (NHJ^SC), 4 + 2N,HA (29)
-------
4. Experiments R«-lating to Process Feasibility
a. Absorption Studies
Initial experiments relating to the Ammonia
Hydrazine Exorption process were concerned with a demonstration that aqueous
hydrazine sulfitc represents a suitable absorbing medium for SO- (Reaction 26),
and with a determination of the extent to which water could be eliminated from
the scrubber system. Inasmuch as hydrazine sulfite cannot be obtained com-
mercially, startup was accomplished with aqueous hydrazine;
(30)
A special apparatus, which was designed and
fabricated for investigating the absorption step, is shown in Figure 42. Also,
two synthetic gas mixtures were specially formulated (by the Ma the son Company]
for use in connection with the apparatus. One of these was a mixture of
0. 3 vol-% SO2 and 99. 7 vol-% NZ. The other conformed to the composition of
a typical flue gas, as shown below, except that water and fly ash were not
included in the formulation, and NO was considered to be NO,
'2'
Flue Gas Composition
at 60° F and 1 atm
i
Component % by Volume
N2 74.9
C02 14.7
H2O 7.25
°2 2'8
S02 0. 3
N0x 0.05
Fly Ash 0. 2 (by weight)
In initial work the SO2~N2 mixture was used for simplicity.
In Figure 42'the gas flows counter-currently
to the liquor in the 1-m. dia. scrubber, designated as 5. Cylinder gas enters
-------
1 -GAS INLET
2-GAS FLOWMETER
3-BUBBLER
4-MANOMETER
5-SCRUBBER
6-HEAT EXCHANGER
7-LIQUID FLOWMETER
8-pH CONTROL VESSEL
9-MIX VESSEL
10-LIQUID MAKEUP
FUNNEL
11-OFFSTREAM
LIQUID SAMPLER
12-LIQUID CIRCULATING
PUMP
13-pH METER
L4-WATER CONDENSER
1 5-GAS SAMPLER SYSTEM
LABORATORY APPARATUS FOR INVESTIGATING S02 REMOVAL FROM FLUE GAS
1 Figure 42
i'
-------
the system at the gas inlet (1) and passes through a flowmeter (2) and a
bubbler (3) (which serves the function of introducing an amount of water
into the gas which is roughly equivalent to that shown in the above table)
before entering the scrubber. The exit gas, of reduced SO2 content, is
cooled by a water condenser (14) to eliminate excess water, and can then
be sampled for. gas chromatographic, mass spectrometric, or other analyses
in any of four gas sample tubes (15). All exiting gas passes through an alkaline
SO, trap (16), the contents of which can be later analyzed for SO2 content.
The counter-current liquid flow from the scrubber
passes through a liquid-seal trap to a mix vessel (9). !which is magnetically
stirred, and into which make-up base (e. g., hydrazine) is added from a dropping
funnel (10). The mixture then flows through a pH control vessel (18), which is
fitted with appropriate electrodes. • The center neck of this flask may contain
a hydrometer (not shown), so that continous specific gravity data may be
obtained. From the pH control vessel the liquid passes through a plastic-lined
circulating pump (12), a heat exchanger (6), which in general maintains the
circulating scrubber liquor at about 50°C, and then back to the scrubber. The
liquid volume of the system, about 500 ml, is maintained at a constant level
through the removal of scrubber liquor to an off stream sampler (11).
i
Calibration runs were made with 20 molal NHL
solutions (20 moles NH-/100 moles H_O) as the absorbent in order to determine
the column absorption efficiency through a comparison with published data. The
data given in Table 73 and in Figure 43 show that the extraction efficiency of
the 1-in. column is considerably higher than that obtained by TVA workers
673
using 8 ft of 2-in. Raschig rings in a 2-ft. dia. column.
The first series of runs made with a. 20 molal
hydrazine solution showed that the extent of SO. absorption was higher than 93%,
*•• A
approximately the same as for the ammonia solutions of comparable pH. The
product gas from Runs H-l and H-2 (see Table 73), made with the 20 molal
hydrazine solution at a 0. 45 SO2/N_H4(S/C) ratio, was analyzed for hydrazine
and ammonia content as well as for SO_, but neither of these compounds was
-------
TABLE 73
Solution
Composition
Run . moles
No. UOO moles Hy
A-l 20 NH3
A-2 "
A-3
A-4
A-5
A-6
A -7 "
H-l 20 N2H4
H-2
H-3 "
H-4 "
H-5
BENCH SCALE
Flow Rates
. Liquid Gas Column A. p
tV (ml/mm) (I/mm) (mm H2O)
178 2.5 18
20
18
" 20
11 20
19
" " 20
178 2.5 17
" " 19
" " 20
" " 22
M n 19
SO2 ABSORPTION DATA
Solution
Temp
50
50
50
50
50
49
50
50
50
50
50
50
Solution
PH
6.55
6.55
6.00
6.00
5.80
5.55
5.55
7.60
7.60
6.80
6.80
6.2
S02m
Feed Gas
(%)
0.279
n
n
n
n
n
M
0.279
ii
ii
n
n
SO2 in
Product Gas
0.020
0.019
0.021
0.021
0.050
0. 155
0. 159
0.018
0.014
0.013
0.018
0.018
S/C Ratio
0.60
0.60
0.675
0.675
0.75
0.82
0.82
0.45
0.45
0.52
0.52
0.60
S = total concentration of dissolved SO, (moles/100 moles water)
C = total concentration of base (moles/100 moles water)
**Determmed by passing effluent gas through standard iodine solution which is subsequently
titrated with standard thiosulfate (SO is quantitatively oxidized by iodine).
so2
Rem
-------
100
90
70
60
~ 50
3 40
•8
1
? 30
«»—
CO
20
10
5.0
a 20 Molal NH3 solution using
27" of 1/4" I ntalox porcelain
saddles .
A20 Molal N2H4 solution using
27" of 1/4" I ntalox porcelain '
saddles
©20 Molal NH3 solution using
8' of 2" Raschig rings
(Data of Hein, Phillips, and -
Young; See Problems and
Control of Air Pollution.
Mallette,F.S.. Editor, New •
York, Reinhold, 1955)
65 .60
8.0
6.0 7.0
Solution pH
ABSORPTION OF S02 BY NH3 AND N2H4 SOLUTION^ IN PACKED COLUMNS
Figure 43
-------
detected. The overall test results indj rated that a 20 molal hydrazine solution
exhibits excellent absorption characteristics for SO, when the S/C ratio is in
the range 0. 45 to 0. 60. For an aqv MUS solution of (N7FI_)_SO, the S/C ratio
£ D b J
is 0. 50, so that it appears from the data that the sulfite is effective in
absorbing SO- in the manner indicated in Reaction 26.
When a 50 molal solution was used for absorption
it was found that a large quantity of white, needle -like crystals precipitated
from the solution when the S/C ratio had increased to a value of 0.45. It was
subsequently noted that no precipitate formed from a 45 molal solution under
similar experimental conditions. An elemental analysis of the solid indicated
that it probably consisted of impure sulfite. The experi-nent w.^s signficant
in indicating the limit to which water could be eliminated from the absorber at
50 C. Although a 50 molal solution (i. e. , C = 50) contains much less water
than that encountered in other scrubber media (e. g. , C < 20 for NHL systems,
and C = 3. 5 for the Johnstone Zinc Oxide process) it appeared that the water
content of the medium was such that objectionable steam costs would still result.
On the other hand the precipitated solid sulfite was readily isolable by filtration,
and it was therefore of interest to consider the possibility of thermally decom-
posing the solid to yield SO- (compare Reaction 30):
(31)
Some encouragement in this direction was afforded by the observation that the
solid melted with gas evolution at the relatively low temperature of 71 C.
' b. Regeneration Studies
(1) Analysis of the Solid
It was noted above that an elemental analysis
of the solid obtained from the N2H4-SO2-H2O system was in reasonable agreement
with that expected for the normal sulfite, (N2H5)2SO3. It was considered, however,
that the solid might be an isomer of fhe sulfite, with the structure
N-H^H-NNHSO ~-H,O, inasmuch as the closely related compound
u o Z 2 Z
-------
(N2H5)2 O2SNHNHSO2 can be isolated from the system
Evidence that the solid was .indeed the sulfite was obtained, however, from
its reaction with aqueous barium chloride. Barium sulfite of 97.5% purity
was precipitated in 89% yield from the reaction mixture. An infrared
of a mulled sample of the original solid also supported the sulfite structure.
An analysis of the solid with aqueous KIO
': ' 3
indicated a purity of 95. 5%, based on the empirical formula N,H,0SO,,
and with all of the nitrogen considered to be in the form of the hydrazine
moiety. With the impurity presumed to be water, on the basis that the
solid was isolated from aqueous solution, the following analytical results
are also definitive: Calc'd, for 95. 5% N4H1QSO3 + 4. 5% HgO; N, 36. 7%;
H, 7.04; S, 20.9. Found: N, 37.6; H, 7.47; S, 20.7.
(2) Thermal Decomposition of Solid Hvdrazine
t Sulfite
In an attempt to effect reaction 31 the
thermal decomposition of the solid was investigated. No appreciable decom-
position was noted below the melting point of the solid at 71 C. At this
temperature gas evolution occurred, leaving a colorless, viscous liquid residue
which did not solidify at room temperature. It was found, however, that the
evolved gas was nitrogen, rather than SO- (by gas chromatography). The
b
quantity of evolved gas accounted for the oxidation of approximately one fourth
of the available hydrazine in the solid. The remainder of the hydrazine was
liberated as such from the solid during the decomposition, and appeared as
a less volatile fraction containing appreciable water. Hydrolysis of the non-
volatile residue with hot HC1 yielded H-S and elemental sulfur as hydrolysis
£*
products, indicating that contained sulfur had undergone reduction during the
thermal decomposition.
(3) Thermal Decomposition of Aqueous Solution
of Hydrazine Sulfite
The reducing properties of aqueous hydrazM
-------
are much mon: pronounced in concentniti-d solution than in dilute solution,
and on this basis it was* considered th it the thermal decomposition of an
aqueous solution of hydrazinc sulfi4 might be less subject to oxidation-
reduction than that attending the decomposition of the solid. This was found
to be the case, but the evolution of nitrogen could not be completely avoided.
For example, when a 40 molal hydrazine solution containing 0. 97 moles of
SO2 per mole of hydrazine was heated to boiling at 70°C and at a pressure
of 20-in. of Hg for 3 hours, about 0. 13 moles of N_ was obtained per mole
of hydrazine. Approximately 35% of the SO. in soluticn in excess of that
required for an S/C ratio of 0. 5 (corresponding to the norm*! sulfite) was
also evolved. In another experiment a 20 molal hydrazine solution with S/C =
0.80 evolved 5. 29% of the theoretical nitrogen during two hours at 67° to 73°C.
The above experiments indicate that the
thermal decomposition of hydrazine sulfite is attended by the formation of
nitrogen under relatively mild conditions. Even a small loss of hydrazine,
a relatively expensive makeup chemical, in the form of elemental nitrogen
is economically prohibitive, and it was therefore decided to terminate work
relating to the hydrazine system.
5. Experiments with Methylhydrazines
I «^—^^
A limited amount of experimentation was carried out
with monomethylhydrazine (MMH) and unsymmetrical dimethylhydrazine
(UDMH) to determine whether these compounds offered any advantage over
hydrazine relative to regeneration of the SO-, as well as of the base. A 20
molal MMH solution was prepared and treated with SO2 to give an S/C ratio
of 0. 7. The solution was found to readily absorb SO2« When heated to the
boiling point (108°C), however, the solution liberated 0.21 moles of nitrogen
gas per mole of MMH over a 75-mm period.
In another experiment a 40 molal solution of UDMH
was prepared and treated with SO2 to give an S/C ratio of unity. When this
solution was subsequently heated SO2> but no nitrogen, was evolved. However
-------
the solution had become intensely yellow, and showed little capacity for the
reabsorption of SO,*
D. THE MITSUBISHI LIME PROCESS
t i
1. Process Description
In the Mitsubishi Lime Process sulfur dioxide is scrubbed
from the flue gas by a 10% hme or limestone slurry. The actual absorbent
is calcium bicarbonate, which is present in solution as the result of an
equilibrium reaction with calcium carbonate. A portion of the circulating
medium is continually withdrawn and oxidized with air to afford a high purity
calcium sulfate (gypsum). The high purity of the product is a consequence of
working with a scrubber effluent which is essentially free of fly ash. A fjlow
diagram of the Mitsubishi Lime process is shown in Figure 44.
2. Process Reactions
Scrubber:
S02+Ca(HC03)2 - » CaS03 J + 2CO2+H2O (32)
• CaS04 { (33)
Makeup Tank:
CaO + CO - » CaCO 1 (34)
2
3 1
and/or
CaC03 ^ H20 + C02 **Ca(HCO3)2 (35)
Oxidation Tower:
CaS03 +1/202 - » CaS04J (33)
3. Process Simplification and Improvement
No experimental work appeared to be indicated at this time
relating to the Mitsubishi Lime process. The process was regarded as eco-
nomical, provided that product gypsum could be sold in quantity. This point
is considered further in Part Four, IV. A.
-------
A
WASTE
GAS COOLER
SCRUBBER NO I
Reactant Chemicals Per Million SCF Flue Gas Processed
Stream
1
2
3
4
Component
so2
S02
Lime ( CaO )
Lb Moles
8.36
0.42
7.94
7.94
Lbs
535
26.8
445
1367
PURIFIED
GAS
SCRUBBER NO 2
"2<>
LIME
LIME MAtw UP TANK
*
AIR
L
GYPSUM
ORYER
OXIDATION
TOWER
CENTRIFUGE
MITSUBISHI LIME PROCESS : FLOW DIAGRAM
-------
PART FOUR
PRELIMINARY PLANT -SCALE PROCESS EVALUATION
COST ESTIMATE - PHASE III
INTRODUCTION
The objective of Phase III was to perform preliminary process designs
and an economic analysis for each candidate process, as applied to the cases
listed below. Data generated in Phases I and II were used as input.
Case Description
1 Large new power
plant facility
1400 megawatt
)
2 Large existing
power plant
facility
1400 megawatt
3 Small existing
power plant
facility
220 megawatt
4 New smelter
facility (5% SO,
to scrubber)
Flue Gas
MMSCFM
2.5
Exit SO-
Plant Factor
Coal
Requirement
tons/hr days/year %-cap.
2.5
150
150
580
580
330
330
100
100
0.5
300
90
330
60
0. 02 5, 000
Not Specified
For Case 4, the extent to which the process would be amenable to fitting into an
existing smelter facility was also of interest.
The flue gas composition for Cases 1-3 was considered to be the same as
that given in Part Two, III. C. 1. a. of this report. The fly ash content (0. 2
wt-%) was presumed to be that which occurs upstream of removal equipment.
In existing plants with installed fly ash collecting equipment it was assumed that
the fly ash content had been reduced to at least 0. 02 weight percent.
An analysis midway in the program indicated that Case 4 as specified
above did not apply to a representative smelter operation; therefore, the initial
-------
conditions were abandoned in favor of a more typical case.
The new conditions were based on data provided to NAPCA in a progress
report by another contractor (Allied Chemical Corp.). The smelter gas in this
case is obtained from a medium size copper smelting facility, and consists of
the combined process gas from three converters and two medium size reverber-
atory furnaces fired with coal. The following conditions prevail:
• Converter cycles are scheduled to even out gas composition, which
varies from 0-21% SO_ during one complete cycle of a converter.
• As per standard American practice, air dilution is used to cool the
converter gas to protect the collecting system.
• Gases from the reverberatory furnaces (100, 000 SCFM) and from
the converters (110,000 SCFM) are combined.
• Based on recent data from an actual operating U.S. copper smelting
facility, converter gas composition is taken as
SO2 - 4.5%
02 - N2 95.5%
and reverberatory gas as
so2 - 1.1%
C02 - 3.5%
CO - 1.2%
H2O - 0.23%
O0 - 13.0%
2
N, - 81.0%
Lt
1.25 Ib/min
L t
Dust - 0. 2 gram/SCF
Combining the gases results in the following standard case:
Smelter Gas - 210,000 SCFM (760 mm & 32°F) or 220,000 SCFM (1 atm
& 60°F - Phase III conditions).
-------
Gas temperature - 440°F
Component % by Volume
S02 2.9
CO2 1.7
CO 0.6
H20 0. 1
02 14.3
N2 80.4
H-SO, 1.251b/mm
i 2 4
Dust 0.2grain/min
It was assumed that the facility would operate at a plant load factor of 90% or
330 days at 100% capacity.
II. PROCESS DESIGN
A. GENERAL
The process design for each case included the operations listed below:
• The quantity and quality or composition of input gas, exit gas,
by-products and "other" raw materials were defined.
• A safety factor of 10% was used for the equipment.
• A preliminary process flow diagram indicating all the operations
required was constructed.
• Each unit operation was checked for alternate routes for possible
economic advantage.
• Alternate equipment routes for SO2 absorption were considered.
These included countercurrent-flow and cross-flow packed
scrubbers, and flooded bed scrubbers (Turbulent Contact Absorber),
the latter for slurry systems.
• Material and energy requirements were prepared.
• The utilities needed were specified.
-------
• Absorption towers and strippers were designed with the
assistance of equipment manufacturers as needed.
• Design data for other equipment, such as pumps, compressors
heat exchangers, filters, centrifugals, crystallizers,
evaporators, agitators, screens, crushers, grinders,
settlers, thickeners, etc. , were obtained from specifications.
The assistance of vendors was requested when necessary.
• A satisfactory material of construction was selected for each
item of equipment.
• Specifications for each item of major process equipment were
prepared.
• A final process flow diagram showing all major items of equipment,
temperatures, pressures, and flow rates in all parts of the process
was prepared. Pertinent process design data were indicated.
Valves, utility lines, and spare items of equipment were omitted.
B. STACK GAS REHEAT
Reheat of the exit gas is needed to effect buoyancy of the gas and also
to prevent a visible plume or condensation of droplets of water from the stack
plvme. From the standpoint of conservation of energy and money, minimum
repeat should be used because tremendous quantities of gases at ambient pressure
must be heated. This involves rather large quantities of energy (gas or low sulfur
oil) if direct heating is used or very large and expensive heat exchange equipment
if indirect heating is used.
Four methods of reheating the exit gas from the aqueous SO, scrubbing
plants are discussed below:
By-passing a part of the hot untreated gas.— This would be feasible
only where less than 90% SO, removal is required. If 100% removal of SO- is
L* Li
achieved in the treated portion, by-passing 10% of the 300°F untreated gas would
gwe an overall SO_ removal of only 90% and would result in a temperature rise
of only about 18°F in the exit gas.
i
Heat exchanging the 300°F inlet gas with 122°F exit gas.— This could
be done in several ways but all are relatively expensive and would be limited to
a reheat of not more than 100°F. One approach is to use a rotating or moving
-------
tf
solid heat sink which would contact first the hot gas and then the cold gas,
transferring the heat to the latter; this is typified by the operation of a boiler
air preheater. In general, this method is used where a temperature differen-
tial, AT, of 100°F or larger, exists.
A second approach would be to pass the two gases counter current
through a heat exchanger having an extended surface on both sides of the tubes
or interfaces. Using this approach for the larger plant (2. 5 MMSCFM feed
gas) results in a .very high heat exchange surface requirement of 758, 000 sq ft.
(This is based on a 110°F reheat, a 44°F mean AT, based on a correction factor
applied to a cross flow shell and 1 tube pass heat exchanger, and a 10 Btu/hr/
sq ft/°F overall transfer coefficient.) An overall transfer coefficient of 10 for
a gas-to-gas exchanger at atmospheric pressure is not conservative unless a
high pressure drop is taken by both fluids. In this case a high pressure drop
would require very expensive gas blowers. At an estimated purchase cost of
$3. 00/sq ft of exchanger surface, the exchanger would cost 2. 3 MM dollars or
9. 1 MM dollars installed (3. 96 factor).
Using a steam coil with extended heat-transfer surface on the gas
side of the exchanger.-- This has the advantage of using high-pressure steam
which would be available at minimum cost at the power plant. The high steam
temperature would give a large AT, thus requiring less surface. For example,
800 psig steam would give a 520°F temperature or a AT of 520-172* = 348°F
(for a 100 reheat); thus, the surface required would be only 44/348 or about
1/8 that required if the inlet gas were exchanged with the exit gas. The dis-
advantages of this system would be the still appreciably high cost of the finned
tube unit and the additional power required by the flue gas blower due to the
additional inch or two of flue gas pressure drop across the finned tube unit.
Direct heating by burners installed in the bottom of the stack. —
This is the method advocated in this report. Either natural gas or low sulfur
oil could be used. The lower section of the stack would have to be lined with
firebrick in order to withstand the higher temperature; however, this cost
would be minor in comparison with the cost of the heat transfer surface
required by indirect heating. The flue gas volume would be increased only about
6% for a 100°F reheat (assuming natural gas burned with 10% excess aj.r).
*A
Average gas temperature
-------
In the initial analyses of Phase III, the gas temperature was increased 50°F
from 122 to 172 F. This was changed toward the end of the effort to an exit
gas temperature of 200 F.
C. SO2 RECOVERY
Recovery of SO2 in marketable forms of liquid SO,, elemental
sulfur and concentrated sulfuric acid was considered. Conversion to sulfuric
acid was the only process used in this study, since reliable data were not
available for the processes involving liquid SO_ or sulfur. Conversion of SO0
ft L,
to SO- was assumed at 96%.
in. ECONOMIC ANALYSIS
A. INTRODUCTION
The economic analysis provided for each case included a capital
cost estimate and an operating cost estimate. The format for capital and
operating cost estimates were changed slightly from that used in Phase I in
order to more nearly conform, when justified, to cost estimates of SO- removal
systems appearing in the recent literature.
B. CAPITAL COST ESTIMATE
Preliminary capital cost estimates were prepared, with total purchased
equipment cost as the basis for a factored cost estimate. Equipment specifications
*
were sent to vendors, who were requested to submit bids. These costs were used
whenever possible. In a few cases it was necessary to supplement these prices
with engineering cost estimates. The capital cost estimate summary form is
shown in Table 74. The total purchased equipment cost, Item 1, was factored
as shown to obtain the other items of fixed capital cost. Table 74 also shows
Working Capital as a part of the total investment. Working Capital, derived from
fixed capital cost and from the operating cost estimate (see below), included the
following items:
&
The companies who assisted in this phase of the cost estimate are acknowledged
in Appendix C -1.
-------
TABLE 74
CAPITAL COST ESTIMATE SUMMARY
ITEM jf) COST - &
1. Purchased Equipment *«0
2. Erection Labor °.25
3. Foundations and Platforms 0. 18 _______
4. Piping ' 0.50
5. Instruments 0. 10
6. Insulation 0.08 __________
7. Electrical 0. 10 '
8. Process Buildings, Structures 0.25 __________
9. Plant Facilities, 5% of 1-8
10. Plant Utilities, 7% of 1-9 ___________
11. PHYSICAL PLANT COST
12. Engineering & Construction, 20% of 11 «_-_-_-^«__i
13. DIRECT PLANT COST
14. Contingency, 15% of 13 _
15. Contractor's Fee, 5% of 13 + 14
16. FIXED CAPITAL COST 4.0
17. Interest During Construction, 2.5% of 16
18. SUB-TOTAL FOR DEPRECIATION
19. Working Capital
20. TOTAL INVESTMENT
21. CAPITAL REQUIREMENTS: $ /kw capacity
-------
• Raw Material inventory - 2 months
• Direct labor - 3 months
• Maintenance - 3 months
• Supplies - 3 months
• Payroll burden - 3 months
• Plant overhead - 4 months
• Fixed cost - 0.5% of fixed capital cost
• Spare parts &
miscellaneous - 1.0% of fixed capital cost
Capital required for plume reheat equipment and for the sulfuric
acid plants were not included as part of the basic plant cost. These were added
as incremental costs since their utility could be applied to other sulfur dioxide
recovery processes.
The fourth method of plume reheat described earlier, i. e. , direct
heating, was used. The capital cost for direct heating is believed to be lower
than for any other system considered.
The capital costs for sulfuric acid plants were obtained by using
data in two literature sources which discussed costs for sulfuric acid plants of
approximately the size required for Cases 1 and 2. ' The requirements for
Cases 3 and 4 were factored from the costs of the larger plant.
C. OPERATING COST ESTIMATE
The operating cost estimates were similar to those developed in
Phase I. Some changes were incorporated, when considered justified, in order
to conform with recent cost estimates of competitive SO2 removal processes
appearing in the literature. Table 75 shows a typical operating cost estimate
summary sheet. Cost elements requiring clarification are discussed below.
• Raw Materials and Chemicals
In Phase I, most of the raw material chemical costs were obtained
from the Oil, Paint and Drug Reporter. In the Phase III work,
chemical manufacturers were contacted for current prices at the
various annual tonnages required.
-------
TABLE 75
OPERATING COST ESTIMATE SUMMARY
ITEM TOTAL $
1. Raw Materials & Chemicals
2. Direct Labor
3. Supervision, 15% of 2
4. Maintenance, 3% of fixed capital cost
5. Supplies', 20% of 4
6. Utilities
7. Other
8. TOTAL DIRECT COST
9. Payroll Burden, IB. 5% of 2 fc 3
10. Plant Overhead, 50% of 2. 3, 4 & 5
11. Waste Disposal
12. Other
13. TOTAL INDIRECT COST
14. Depreciation. 11% depreciable capital cost
15. Taxes & Insurance, 3% depreciable cap. cost
16. Other
17. TOTAL FIXED COST
18. TOTAL OPERATING COST 100.0_
19. Cost: $ /ton coal. mill/kwh
-------
• Direct Labor
The hourly wage rate was $2. 75 per hour.
• Utilities
Utilities were charged at the same unit costs as those used in
Phase I, as follows:
Steam, M Ib $0. 50
Heat credits &c debits, MM Btu 0. 50
Power, kwh 0.006
Raw water, M gal 0. 10
Recirculated water, M gal 0. 05
Fuel oil, gallon 0. 10
• Waste Disposal
Costs of disposal of waste materials were applied when applicable.
• Total Operating Cost
The total operating cost did not include any credit for by-product
sales.
D. PROFITABILITY
The profitability was checked for each system in which the process
involved a salable by-product. This was presented in both tabular form and as
a break-even chart indicating by-product sales at various price levels.
IV. PROCESSES EVALUATED
A. INTRODUCTION
As mentioned in Part One, four processes of the twenty-two checked
in Phase I were selected as candidates for additional effort in Phases II and III.
Two of these, the Ammonia-Hydrazine Exorption and Cominco Exorption Processes
were eliminated from further contention in Phase III due to the negative results
of the laboratory experimentation, Phase II. This left two processes for Phase III
evaluation: the Zinc Oxide Process and the Mitsubishi Lime Process.
The Phase III evaluation of Cases 1, 2, 3, and 4 of the Zinc Oxide
-------
Process are reported in detail in this section.
A re-evaluation of the potential gypsum market in the United States
led to the conclusion that it may be more realistic to consider a simplified Ume
process in the Phase III effort rather than the Mitsubishi Lime process. In the
former case, gypsum would be disposed of as a waste material along with the
fly ash rather than recovered in pure form for marketability as is done in the
Mitsubishi Lime process. Accordingly, the Howden-I. C. I. or the essentially
equivalent Mitsubishi Simplified Lime process would form the basis for the
Phase ILL study.
Information acquired in a recent conversation with the president of
the Southern California Gypsum Association illustrates the dim forecast for
promoting the marketability of pure gypsum as obtained in the Mitsubishi Lime
process. Natural deposits of gypsum are very widespread in this country and
manufacturers using gypsum for wallboard, etc., have acquired large reserves
which will last for 20-50 years. Moreover, these manufacturers have sub-
stantial investments in mining and transportation equipment needed to exploit
these reserves. -Accordingly, it is difficult to estimate the cost of gypsum to
the manufacture!;, or more important how much a manufacturer would be willing
to pay for gypsum from a source such as a power plant. Except for isolated
cases where the power plant producing gypsum is located reasonably close to a
gypsum manufacturer who would normally have to receive the material from
a distant source, the selling price of the gypsum would probably be only a
dollar or two a ton. Further, the acceptance of "power-plant" gypsum would
be questionable, because although it may be purer than the natural product, its
hardness, size, and processing qualities would be different and the manufacturing
processes and/or equipment would probably have to be modified to use it. For
these reasons it was decided that the lime process would be treated as a simpli-
fied system which involved no product recovery. This plan was not carried to
completion, however, since other NAPCA investigators were giving adequate
study to processes utilizing lime and limestone. A Case 3 analysis for the
Simplified Lime Process was conducted, however, which is presented later.
-------
B. ZINC OXIDE PROCESS
1. Process Design
The same basic process design was used for all cases of the
Zinc Oxide Process. This consisted of the original work by Johnstone, et al.
and on data developed by the Bureau of Mines. A departure from the design
used in Cases 1, 2, and 3 was made in Case 4 which involved dust removal. A
prescrubber has been substituted in Case 4 for the dry cyclone used previously.
The prescrubber selected was assumed to absorb all of the H-SO. and remove
95% of the dust in the smelter gas. Material balances and flow diagrams were
prepared. All major equipment was sized and specifications were written. The
process flow diagrams indicate the entire processing sequence and include
equipment sizing information, flows, temperatures, and pressures. Tentative
dimensions are given for pipe diameters, in inches, and ductwork, in feet.
These dimensions are based on liquid flows of 5. 5 ft/sec or less and gas flows
of 50 ft/sec or less. Raw material and utility requirements were estimated.
The process flow diagrams are shown in Figures 45-47.
2. Capital Costs
The capital cost estimates provided the following investment
requirements:
Case Total Investment, $ $/kw Capacity
1 15,688,700 11.21
2 12,435,600 8.88
3 3,510,000 15.95
4 9,213,500
Capital cost estimate summaries are presented for each case
(see Tables 76-79). Equipment lists showing basic data and individual equipment
estimated costs are ID Appendix C. The equipment list for Cases 1 and 2 is
identical except that the flue gas cyclone is considered to be a part of existing
power plant equipment in Case 2. Derivation of working capital is shown in
Tables 80-83.
-------
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-------
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-------
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-------
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-------
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Figure 46
-------
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-------
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-------
TABLE 76
ZINC OXIDE PROCESS: CAPITAL COST ESTIMATE SUMMARY
Case 1
ITEM COST - $
1. Purchased Equipment 3,675,000
2. Erection Labor 918,800
3. Foundations and Platforms 661,500
4. Piping 1,837,500
5. Instruments 367,500
6. Insulation 294,000
7. Electrical 367,500
8. Process Buildings, Structures 918.800
9. Plant Facilities, 5% of 1-8 452,000
10. Plant Utilities, 7% of 1-9 664,500
11. PHYSICAL PLANT COST 10,157,100
12. Engineering & Construction, 20% of 11 2,031,400
13. DIRECT PLANT COST 12,188,500
14. Contingency, 15% of 13 1,828,300
15. Contractor's Fee, 5% of 13 + 14 700.800
16. FIXED CAPITAL COST 14, 717.600
17. Interest During Construction, 2. 5% of 16 ^«^iZi222.i
18. SUB-TOTAL FOR DEPRECIATION 15,085,500
19. Working Capital 603,200
20. TOTAL INVESTMENT 15, 688. 700
21. CAPITAL REQUIREMENTS: $ 11.21 /kw capacity
-------
TABLE 78
ZINC OXIDE PROCESS: CAPITAL COST ESTIMATE SUMMARY
Case 3
ITEM COST - $
1. Purchased Equipment 815,000
2. Erection Labor 203,800
3. Foundations and Platforms 146, 700
4. Piping 407,500
5. Instruments 81,500
6. Insulation 65,200
7. Electrical 81,500
8. Process Buildings, Structures 203,800
9. Plant Facilities, 5% of 1-8 1QQ, 300
10. Plant Utilities, 7% of 1-9 147.400
11. PHYSICAL PLANT COST 2.252,700
12. Engineering & Construction, 20% of 11 450,500
13. DIRECT PLANT COST 2. 703.200
14. Contingency, 15% of 13 405.500
15. Contractor's Fee, 5% of 13 + 14 155.400
16. FDCED CAPITAL COST 3.264. 100
17. Interest During Construction, 2. 5% of 16 81.600
18. SUB-TOTAL FOR DEPRECIATION 3. 345. 700
19. Working Capital *68, 50°
20. TOTAL INVESTMENT 3. 514. 200
i
21. CAPITAL REQUIREMENTS: $ 15.97 /kw capacity
-------
TABLE 79
ZINC OXIDE PROCESS: CAPITAL COST ESTIMATE SUMMARY
Case 4
ITEM CQST-$
1. Purchased Equipment • 2, 140,000
2. Erection Labor 535,000
3. Foundations and Platforms 385.200
4. Piping 1.070,000
5. Instruments 214.000
6. Insulation 171.200
7. Electrical 214.000
8. Process Buildings, Structures 535, OOP
9. Plant Facilities, 5% of 1-8 263.200
10. Plant Utilities, 7% of 1-9 386.900
11. PHYSICAL PLANT COST 5.914.500
12. Engineering & Construction, 20% of 11 1. 182. 900
13. DIRECT PLANT COST 7.097.400
14. Contingency, 15% of 13 1.064.600
15. Contractor's Fee, 5% of 13 + 14 ^408.100
16. FIXED CAPITAL COST 8.570.100
17. Interest During Construction, 2. 5% of 16 214, 300
18. SUB-TOTAL FOR DEPRECIATION 8.784.400
19. Working Capital 460, OOP
20. TOTAHNVESTMENT 9.213.500
-------
TABLE 80
ZINC OXIDE PROCESS: WORKING CAPITAL
Case 1
70% Plant Factor
COST - $
Raw Material Inventory, 2 months
Direct Labor, 3 months
Maintenance, 3 months
Supplies, 3 months
Payroll Burden, 3 months
Plant Overhead, 4 months
Fixed Cost, 0. 5% fixed capital cost
Spare Parts & Miscellaneous,
1.0% fixed capital cost
Sub-total:
ZnO
98, 100
32, 100
110,400
22, 100
6,800
112,900
73,600
147,200
603,200
H2S04
-
24, 700
50,000
10,000
5,300
58, 900
20,000
40,000
208, 900
Plume
Reheat
-
-
1,500
300
-
1,200
1,000
2,000
6,000
Total:
$818,100
-------
TABLE 81
ZINC OXIDE PROCESS: WORKING CAPITAL
Case 2
70% Plant Factor
COST - $
Raw .Material Inventory, 2 months
Direct Labor, 3 months
Maintenance, 3 months
Supplies, 3 months
Payroll Burden, 3 months
Plant Overhead, 4 months
Fixed Cost, 0. 5% fixed capital cost
Spare Parts & Miscellaneous,
1. 0% fixed capital cost
Sub-total:
ZnO
98, 100
32, 100
87,300
17,500
6,800
94, 400
58,200
116,300
H2S04
-
24, 700
50,000
10,000
5,300
58, 900
20,000
40,000
Plume
Reheat
-
-
1,500
300
-
1,200
1,000
2,000
510,700 208,900 6, OOC
Total:
$725,600
-------
TABLE 82
ZINC OXIDE PROCESS: WORKING CAPITAL
Case 3
70% Plant Factor
COST - $
Raw Material Inventory, 2 months
Direct Labor, 3 months
Maintenance, 3 months
Supplies, 3 months
Payroll Burden, 3 months
Plant Overhead, 4 months
Fixed Cost, 0. 5% fixed capital cost
Spare Parts & Miscellaneous,
1.0% fixed capital cost
Sub-total:
ZnO
18,900
26, 100
24, 500
4,900
5,600
39,600
16,300
32,600
168,500
H2S°4
-
24, 700
16,500
3,300
5,300
32, 100
6,600
13,200
101,700
Plume
Reheat
-
500
100
-
400
400
700
2, 100
Total:
$272,300
-------
TABLE 83
ZINC OXIDE PROCESS: WORKING CAPITAL
Case 4
Plant Factor
Raw Material Inventory, 2 months
Direct Labor, 3 months
Maintenance, 3 months
Supplies, 3 months
Payroll Burden, 3 months
Plant Overhead, 4 months
Fixed Cost, 0. 5% fixed capital cost
Spare Parts & Miscellaneous,
1.0% fixed capital cost
Sub-total
TOTAL:
COST - $
ZnO
139,200
32, 100
64, 300
12, 900
6,800
76, 100
42, 900
85, 700
Sz§°4
_
$24,700
45,000
9,000
6,800
54, 900
18,000
36, 000
flume
Reheat
«.
-
$400
100
-
300
200
500
460,000 194,400 1,500
$655,900
-------
3. Operating Costs
The initial economic analyses of the cases evaluated in Phase III
assumed a 90% plant factor (330 days operation per year). In Cases 1, 2 and 4
operations were considered to be continuous under these conditions. In Case 3,
however, operations averaged 60% of capacity on the operating days, equivalent
to a 54% plant factor. In all cases a plume reheat of 50°F (a rise from 122° to
172°F) was used.
Some of the guidelines for operations of SO, recovery systems
jjC "
in conjunction with power plants were changed by NAPCA after this work was
completed. It was decided that the 90% plant factor was much too high for most
power plant operations due to the wide fluctuations in power load created by the
variable demand for power. A decision was made to use a 70% plant factor in
these cases. It was also decided to standardize on the plume reheat temperature.
This is now set at a plume temperature of 200°F minimum and at least 50 F above
the exit gas dewpoint.
The estimated operating costs in accordance with the new
guidelines are presented below. These costs do not include by-product credits.
The effect of sulfunc acid sales on these costs is discussed in the next section.
Case 1
System M$/year $/ton coal mill/kwh
5,082
502
1, 745
7,329
Case 2
4,473
502
1.745
6,720
1.43
0. 14
0.49
2.06
1.26
0. 14
0.49
1.89
0.59
0.06
0.20
0.85
0.52
0.06
0.20
0. 78
ZnO Process
Plume Reheat
Sulfuric Acid
Total
ZnO Process
Plume Reheat
Sulfuric Acid
Total
NAPCA Contractors' Meeting, Cincinnati, Ohio on December 12, 1968.
-------
System M$/year $/ton coal mill/kwh
ZnO Process 1,225 2.22 0.91
Plume Reheat 106 0. 19 0.08
Sulfuric Acid 618 1. 12 0.46
Total 1,949 3.53 1.45
Case 4
ZnO Process
Plume Reheat
Sulfuric Acid
Total 5,734 - -
A 90% plant factor was used in Case 4 since the operations of
a smelter facility would not be subject to the load fluctuations experienced in
power plant operations.
Operating cost summary sheets are presented for Cases 1-4
according to the new guidelines (in Tables 84-87. Individual operating cost data
sheets are presented for the plume reheat facility and sulfuric acid plants as
Tables 88-93. Detail sheets for each case show the raw material requirements
and costs, manning table and costs, and utility requirements and costs. These
data are presented in Tables 94-102.
Raw material requirements were obtained from material
balance data shown on the flow diagrams. Raw material purities have been
accounted for in the estimate of their usage. In-plant loss allowances were
selected on the basis of total usage, i. e. , 0. 2% for solid materials (ZnO and
lime) and 0. 1% of the absorbing liquid (Na_O equivalent). The raw material
requirement tables indicate a substantial reduction in raw material cost if
limestone could be substituted for lime. The use of limestone apparently has
not been tried in this process.
The manning table shows the direct labor requirement for
each cost center.
-------
TABLE 84
ZINC OXIDE PROCESS: OPERATING COST ESTIMATE SUMMARY
Case 1
70% Plant Factor
Fixed Capital Cost: $14, 717, 600
ITEM TOTAL $
1. Raw Materials k Chemicals 588,400 11.58
2. Direct Labor 128.500 2.53
3. Supervision, 15% of 2 ig. 300 Q. 38
4. Maintenance, 3% of fixed capital cost 441. 500 8. 69
5. Supplies, 20% of 4 88.300 1. 74
6. Utilities 1.277.000 25. 12
7. Other
8. TOTAL DIRECT COST 2. 543.000 50.04
9. Payroll Burden, 18. 5% of 2 & 3 27. 300 0.54
10. Plant Overhead, 50% of 2, 3, 4 & 5 338.800 6. 67
11. Waste Disposal 61. 100 1.20
12. Other ^—^mm-. __^__
13. TOTAL INDIRECT COST 427.200 8.41
14. Depreciation, 11% depreciable capital cost 1.659.400 32.65
15. Taxes & Insurance, 3% depreciable cap. cost 452. 600 8. 90
16. Other - -
17. TOTAL FDCED COST 2. 112.000 41.55
18. TOTAL OPERATING COST 5.082.200 100*0
19. Cost: $ 1.43 /ton coal, 0.59 mill/kwh
-------
TABLE 85
ZINC OXIDE PROCESS: OPERATING COST ESTIMATE SUMMARY
Case 2
70% Plant Factor
Fixed Capital Cost: $11, 634, 000
ITEM
1. Raw Materials fc Chemicals
2. Direct Labor
3. Supervision, 15% of 2
4. Maintenance, 3% of fixed capital cost
5. Supplies1, 20% of 4
6. Utilitiea
7. Other
8. TOTAL DIRECT COST
9. Payroll Burden, 18. 5% of 2 & 3
10. Plant Overhead, 50% of 2, 3, 4 fc 5
11. Waste Disposal
12. Other
13. TOTAL INDIRECT COST
TOTAL $
588,400
128.500
19.300
349.000
69. 800
1.277.000
2.432.000
27.300
283.300
61.100
371.700
14. Depreciation, 11% depreciable capital cost 1. 311. 700
15. Taxes fc Insurance, 3% depreciable cap. cost 357. 700
16. Other _
17. TOTAL FDCED COST ,
18. TOTAL OPERATING COST
19. Cost: $ 1.26 /ton coal, 0.52 mill/kwh
4.473. 100
0.44
7.80
1.56
28.55
54.37
0.61
1.37
29.32
-------
TABLE 86
ZINC OXIDE PROCESS: OPERATING COST ESTIMATE SUMMARY
Case 3
70% Plant Factor
Fixed Capital Cost: $3, 264, 100
ITEM TOTAL $ %
1. Raw Materials It Chemicals 113.200 9.24
2. Direct Labor 104.400 8.52
3. Supervision. 15% of 2 15. 700 \t 28
4. Maintenance, 3% of fixed capital cost 97i 900 7.99
5. Supplies, 20% of 4 19.600 1.60
6. Utilities 248.300 20.26
7. Other
8. TOTAL DIRECT COST 599. 100 48.89
9. Payroll Burden, 18.5% of 2 & 3 22.200 1.81
10. Plant Overhead, 50% of 2, 3, 4 & 5 us. 800 9. 70
11. Waste Disposal 16.800 l. 37
12. Other . .
13. TOTAL INDIRECT COST 157.800 12.88
14. Depreciation, 11% depreciable capital cost 368. OOP 30. 03
15. Taxes & Insurance, 3% depreciable cap. cost 100.400 8. 20
16. Other - -
17. TOTAL FIXED COST 468.4QQ 38.23
18. TOTAL OPERATING COST 1.225,300 10°-°
19. Cost: $ 2.22 /ton coal, 0.91 mill/kwh
-------
TABLE 87
ZINC OXIDE PROCESS: OPERATING COST ESTIMATE SUMMARY
Case 4
90% Plant Factor
\ ,
Fixed Capital Cost: $8, 570, 100
ITEM TOTAL $
1. Raw Materials fc Chemicals 649.7QQ 16. a&
2. Direct Labor 128.500 3.24
3. Supervision, 15% of 2 19.300 0.49
4. Maintenance, 3% of fixed capital cost 257.100 6.48
5. Supplies, 20% of 4 51.400 1.29
6. Utilities. 1.319.000 33.22
7. Other . -
8. TOTAL DIRECT COST , 2,425.000 61.08
9. Payroll Burden, 18. 5% of 2 fc 3 27. 300 0.69
10. Plant Overhead, 50% of 2, 3, 4 & 5 228.200 5.75
11. Waste Disposal 60. OOP 1.51
12. Other - _^L«.
13. TOTAL INDIRECT COST 315.500 7.95
14. Depreciation, 11% depreciable capital cost 966. 300 24.34
15. Taxes fc Insurance, 3% depreciable cap. cost 263.500 6.64
16. Other . .
17. TOTAL FIXED COST 1.229.800 30.97
18. TOTAL OPERATING COST 3,970.300 100.0
-------
TABLE 88
PLUME REHEAT: OPERATING COST ESTIMATE SUMMARY
122-200°F
Cases 1 & 2
70% Plant Factor
Fixed Capital Cost: $200, 000
ITEM TOTAL $
1. Raw Materials & Chemicals - -
2. Direct Labor - -
3. Supervision, 15% of 2 - -
4. Maintenance, 3% of fixed capital cost 6.000 1.20
5. Supplies, 20% of 4 1.2QQ 0.24
6. Utilities 462.000 92. 12
7. Other
8. TOTAL DIRECT COST 469.200 93.56
9. Payroll Burden, 18. 5% of 2 & 3 - -
10. Plant Overhead, 50% of 2, 3, 4 & 5 3.600 0. 72
11. Waste Disposal - -
12. Other
13. TOTAL INDIRECT COST 3.600 Q. 72
14. Depreciation, 11% depreciable capital cost 22.600 4.50
15. Taxes fc Insurance, 3% depreciable cap. cost 6. 100 1-22
16. Other - -
17. TOTAL FIXED COST 28. 700 5.72
18. TOTAL OPERATING COST 501.500 100.0
19. Cost: $ 0. 141 /ton coal, 0.058 mill/kwh
-------
TABLE 89
PLUME REHEAT: OPERATING COST ESTIMATE SUMMARY
122-20QPF
Case 3
70% Plant Factor
Fixed Capital Cost: $70, 000.
ITEM TOTAL $
1. Raw Materials & Chemicals - -
2. Direct Labor - -
3. Supervision, 15% of 2 - _.
4. Maintenance, 3% of fixed capital cost 2. 100 1. 98
5. Supplies, 20% of 4 400 0.38
6. Utilities 92.400 86.92
7. Other - -
s. TOTAL DIRECT COST 94.900 39.28
9. Payroll Burden, 18. 5% of 2 fc 3 - -
10. Plant Overhead, 50% of 2, 3, 4 fc 5 1.300 1.22
11. Waste Disposal - -
12. Other - -
13. TOTAL INDIRECT COST 1.300 1.22
14, Depreciation, 11% depreciable capital cost 7. 9QQ 7.43 _
15. Taxes & Insurance, 3% depreciable cap. cost 2. 200 2.07 __
16. Other . .
17. TOTAL FDCED COST 1Q. 1QQ 9.50
18. TOTAL OPERATING COST 106,300 100.0 _
19. Cost: $ 0. 19 /ton coal, 0.079 mill/kwh
-------
TABLE 90
PLUME REHEAT: OPERATING COST ESTIMATE SUMMARY
122-200°F
Case 4
90% Plant Factor
Fixed Capital Cost: $45, 000
ITEM TOTAL $ %
1. Raw Materials & Chemicals - -
2. Direct Labor - -
3. Supervision. 15% of 2 - -
4. Maintenance, 3% of fixed capital cost 1.400 2. 79
5. Supplies, 20% of 4 300 0.60
6. Utilities 41.000 81.84
7. Other
8. TOTAL DIRECT COST 42.700 85.23
9. Payroll Burden. 18. 5% of 2 & 3 - -
10. Plant Overhead, 50% of 2, 3, 4 & 5 900 1.80
11. Waste Disposal - -
12. Other - -
13. TOTAL INDIRECT COST 900 i.flO
14. Depreciation, 11% depreciable capital cost 5, 100 10. 18
15. Taxes fc Insurance, 3% depreciable cap. cost 1.400 2.79
16. Other - -
17. TOTAL FDCED COST 6.500 12.97
18. TOTAL OPERATING COST 50. 100 100-0
-------
TABLE 91
SULFURIC ACID PLANT: OPERATING COST ESTIMATE SUMMARY
Cases 1 & Z
,70% Plant Factor
Fixed Capital Cost: $4, 000, 000
ITEM TOTAL $
1. Raw Materials fc Chemicals -0-
2. Direct Labor 98.600 5.65
3. Supervision, 15% of 2 14.800 0.85
4. Maintenance, 3% of fixed capital cost 200.000 11.46
5. Supplies, 20% of 4 40.000 2.29
6. Utilities 619.500 35.52
7. Other - -
8. TOTAL DIRECT COST 972.900 55.77
9. Payroll Burden, 18. 5% of 2 It 3 21.000 1.20
10. Plant Overhead, 50% of 2, 3, 4 & 5 176. 7QQ 10. 13
11. Waste Disposal • -
12. Other - - .
13. TOTAL INDIRECT COST 197.700 11.33
14. Depreciation, 11% depreciable capital cost 451. OOP 25.85
15. Taxes & Insurance, 3% depreciable cap. cost 123. OOP __ 7.05
16. Other . - .
17. TOTAL FKED COST 574.000 32.9Q
18. TOTAL OPERATING COST 1, 744,600 100.0_
19. Cost: $ 0.49 /ton coal, p. 20 mill/kwh. $5. 68/ton H;.SO4
-------
TABLE 92
SULFURIC ACID PLANT: OPERATING COST ESTIMATE SUMMARY
Case 3
70% Plant Factor
Fixed Capital Cost: $1, 320, 000
ITEM
1. Raw Materials & Chemicals
2. Direct Labor
3. Supervision, 15% of 2
4. Maintenance, 3% of fixed capital cost
5. Supplies, 20% of 4
6. Utilities
7. Other
8. TOTAL DIRECT COST
9. Payroll Burden, 18. 5% of 2 & 3
10. Plant Overhead, 50% of 2, 3, 4 & 5
11. Waste Disposal
12. Other
13. TOTAL INDIRECT COST
18. TOTAL OPERATING COST
TOTAL $
-0-
98.600
14.800
66.000
13.200
118.500
311. 1QQ
21.000
96.300
117.300
14. Depreciation, 11% depreciable capital cost 148. 800
15. Taxes & Insurance, 3% depreciable cap. cost 40.600
16. Other -
17. TOTAL FDCED COST 189.400
617.800
15.96
2.40
10.68
2. 14
24. 78
50.36
3,40
15.59
18.99
24.09
6.56
30.65
100.0
19. Cost: $ l. 12 /ton coal. p. 46 mill/kwh. $10. 71/ton H2SO4 (100%)
-------
TABLE 93
SULFURIC ACID PLANT: OPERATING COST ESTIMATE SUMMARY
Case 4
90% Plant Factor
Fixed Capital Cost: $3, 600, 000
ITEM TOTAL$
1. Raw Materials & Chemicals - ____-__
2. Direct Labor 98.600 5.75
3. Supervision, 15% of 2 14.800 0.86
4. Maintenance, 3% of fixed capital cost 180. OOP 10.52
5. Supplies,, 20% of 4 36. OOP 2.10
6. Utilities 681.8PP 39.79
7. Other
8. TOTAL DIRECT COST 1.011.200 59.02
9. Payroll Burden, 18.5% of 2 fc 3 21.000 1.22
10. Plant Overhead, 50% of 2, 3, 4 It 5 164. 700 9.61
11. Waste Disposal - . -
12. Other . .
•••••••••••• ••••••••
13. TOTAL INDIRECT COST 185.700 10.83
14. Depreciation, 11% depreciable capital cost 405, 900 ' 23.69
15. Taxes fc Insurance, 3% depreciable cap. coat 110.700 6.46
16. Other - -
17. TOTAL FIXED COST 516.600 30.15
18. TOTAL OPERATING COST ;1,713,500 100.0
-------
TABLE 94
ZINC OXIDE PROCESS:
ANNUAL RAW MATERIAL REQUIREMENTS AND COSTS
Cases 1 & 2
70% Plant Factor
Material
Zinc Oxide
Soda Ash
Lime
Cost
$ /unit
0. 15/lb
32. 70 /ton
18 /ton
Quantity
tons /year
626
1,518
19,500
Total:
Total Cost
$/year
187,800
49, 600
351,000
588,400
If limestone could be substituted for lime:
Limestone 2. 50/ton* 35,500 88,800
Total: 326,200
Includes: $1. 35/ton - f. o. b. mines
0.70/ton - freight
0. 37/ton - grinding costs
-------
TABLE 95
ZINC OXIDE PROCESS:
ANNUAL RAW MATERIAL REQUIREMENTS AND COSTS
Case 3
70% Plant Factor
Material
Zinc Oxide
Soda Ash
Lime
Cost
$/unit
0. 15/lb
32. 70 /ton
18 /ton
Quantity
tons /year
118
333
3,716
Total Cost
$/year
35,400
10,900
66,900
Total: 113,200
If limestone could be substituted for lime:
Limestone 2.50/ton 6,670 16,700
Total: 63,000
Includes: $1. 35/ton - f. o. b. mines
0. 70/ton - freight
0. 37/ton - grinding costs
-------
TABLE 96
ZINC OXIDE PROCESS:
ANNUAL RAW MATERIAL REQUIREMENTS AND COSTS
Case 4
90% Plant Factor
Material
Zinc Oxide
Soda Ash
Lime
Cost
$/ unit
0. 15/lb
32. 70 /ton
18 /ton
Quantity
tons /year
696
1673
21,455
Total
Total Cost
$/year
208, 800
54, 700
386, 200
649, 700
If limestone could be substituted for lime:
Limestone 2.50/ton* 39,025 97.600
Total 361,100
Includes: $1. 35/ton - f. o. b. mines
0. 70/ton - freight
0. 37/ton - grinding costs
-------
TABLE 97
ZINC OXIDE PROCESS:
MANNING TABLE AND COST
Cases 1 & 2
Manhours/day-
Shift No.
Operation _1_ _2_ _3_ Total
Absorption 888 24
Regeneration 888 24
i
Drying and Calcining 888 24
Desulfation 888 24
Waste Handling 888 24
Raw Materials 8 8
Total, manhours 48 40 40 128
Total, men 655 16
Cost; 128 hr/day x 365 day/yr x $2. 75/hr = $128, 500
-------
TABLE 98
ZINC OXIDE PROCESS:
MANNING TABLE AND COST
Case 3
Manhours/day
Operation
Absorption
Regeneration & Desulfation
Drying & Calcining
Waste Handling
Raw Materials
Total, manhours
Total, men
Shift No.
J_
8
8
8
8
8
40
5
_2_
8
8
8
8
-
32
4
_3_
8
8
8
8
-
32
4
Total
24
24
24
24
8
104
13
Cost: 104 hr/day x 365 day/yr x $2. 75/hr = $104, 400
-------
TABLE 99
ZINC OXIDE PROCESS:
MANNING TABLE AND COST
Case 4
Operation Manhours /day
Shift No.
Total
Absorption 888 24
Regeneration 888 24
Drying and Calcining 888 24
Desulfation 888 24
Waste Handling 888 24
Raw Materials 8 - - 8
Total, manhours 48 40 40 128
Total, men 655 16
Cost: 128 hr/day x 365 day/yr x $2. 75/hr * $128, 500
-------
TABLE 100
ZINC OXIDE PROCESS:
ANNUAL UTILITY REQUIREMENTS AND COSTS
Cases 1 & 2
70% Plant Factor
Utility
Power
Raw Water
Recirculated Water
Fuel Oil
Cost
$/umt
0.006/kwh
0. 10/M gal
0.05/M gal
2. 92/bbl
Quantity
units /year
32,400,000
336,000
1,355,000
336,000
Total Cost
$/year
194, 500
33,600
67,800
981, 100
Total:
$1,277,000
-------
TABLE 101
ZINC OXIDE PROCESS:
.ANNUAL UTILITY REQUIREMENTS AND COSTS
Case 3
70% Plant Factor
Utility
Power
Raw Water
Recirculated Water
Fuel Oil
Cost
$/unit
0.006/kwh
0. 10/Mgal
0.05/Mgal
2.92/bbl
Quantity
units /year
6,888,600
74, 300
281,200
63,500
Total Cost
$/year
41,400
7,400
14, 100
185,400
Total:
$248,300
-------
TABLE 102
ZINC OXIDE PROCESS:
ANNUAL UTILITY REQUIREMENTS AND COSTS
Case 4
90% Plant Factor
Utility
Power
Raw Water
Re circulated
Water
Fuel Oil
Cost
$/Unit
0. 006/kwh
0. 10/Mgal
0.05/Mgal
2.92/bbl
Quantity
Unite /Year
24, 800, 000
136,000
1,500,000
370, 000
Total:
Total Cost
$/Year
149, 000
14, 000
75, 000
1,081,000
1,319,000
-------
Power consumption was estimated from the total equipment
horsepower specifications. Raw water needs were obtained from make-up
requirements and estimated losses. The recirculated water is based on the
cooling loads specified. Fuel oil is indicated as the source of heat for the
calciner and dryer, although natural gas could be used as an alternate fuel.
4. Profitability
A summary of the economic analysis of each case is presented
below. It was assumed the sulfuric acid would be sold with credits applied
against operating costs. Various prices are shown for sulfuric acid. The
$34/ton indicates the approximate selling price for 100% sulfuric acid in
February 1968. The $23. 50/ton value is an arbitrary selling price equivalent
to about 70% of that price. The other sulfuric acid prices shown indicate the
lowest prices at which the operations would break-even at full capacity* i.e.,
at 70% plant factor for Cases 1-3, and 90% plant factor for Case 4.
The data show that the Zinc Oxide Process may be economically
sound under certain favorable marketing conditions when the SO. is converted
£t
and sold as concentrated sulfuric acid; whether the acid marketing situation
will ever be so favorable as to permit operation of this SO, control process at
£
near break-even conditions is doubtful.
Details of the economic analyses are presented in Tables 103-
106. The data are shown as break-even charts in Figures 48-51.
Case 1
Sulfuric Acid Sales, $/ton
34 24.85 23. 50
Profit Before Tax
$/ton of coal Q. 76 -0- -0-
mills/kwh o.31 -0- -0-
$/tonofH2S04 9>15 ,0. _Q_
Break-even point 43% 70% 73%
-------
Case 2
Sulfunc Acid Sales, $/ton
34 24.85 23.50
Profit Before Tax
$/tonofcoal 0.93 0.06 -0-
mills/kwh 0.38 0.02 -0-
$/ton of H2SO4 11.22 1.61 -0-
Break-even point 37% 66% 70%
The economics of Case 2 would be at approximately the same
level as in Case 1.
Case 3
Sulfuric Acid Sales, $/ton
34 23.50
Profit (loss) Before Tax
$/ton of coal (0. 12) (1. 19)
mills/kwh (0.05) (0.50)
$/tonofH2S04 (1.17) (11.65)
Break-even point 73% —
In Case 3, operation of the SO- removal system would result
in a loss, even if acid could be sold at these very high prices.
Case 4
Sulfuric Acid Sales, $/ton
34 23.50 17. 70
Profit Before Tax
$/ton of H2SO4 16.30 5.80 -0-
Break-even Point 33% 55% 90%
The economics of Case 4 would be more attractive than the
others due to the relatively large quantity of recovered
-------
TABLE 103
ZINC OXIDE PROCESS: ECONOMIC ANALYSIS
i
Case 1
70% Plant Factor b 96% Conversion of SO2 to SO3
I. Capital Investment
Zinc Oxide Process $15, 085, 500
Plume Reheat 205, 000
Sulfunc Acid 4, 100, 000
Working Capital 818, 100
Total Investment: ~ $20, 209, 000
II. Profitability
Sulfuric Acid, $/ton
^$34 $24.85 $23.50
Sales, M $, 295 M tons H2SO4 10,030 7,329 6,930
Operating Cost, M $
Zinc Oxide Process 5,082 5,082 5,082
Plume Reheat 502 502 502
Sulfuric Acid 1,745 1.745 1.745
Total Operating Cost: 7,329 7,329 7.329
Profit Before Tax, M$ 2,701 -0- (399)
Profit After Tax, M$ 1,296 -0- -0-
Return on Total Investment A/T, % 6.4 -0- -0-
Payout, years 5. 9 -0-
Profit Before Tax
$/ton of coal 0.76 -0- -0-
mills/kwh 0.31 -0- -0-
$/tonofH2SO4 9.15 -0- -0-
Break-even Point 43% 70%
-------
TABLE 104
ZINC OXIDE PROCESS: ECONOMIC ANALYSIS
Case 2
70% Plant Factor b 96% Conversion of SO, to SO-
2 3
I. Capital Investment
Zinc Oxide Process $11,924,900
Plume Reheat 205, 000
Sulfunc Acid 4, 100, 000
Working Capital 725,600
Total Investment: ~ $16,956,000
II. Profitability
Sulfuric Acid, $/ton
$23.50 $22. 78
Sales, M$, 295 M tons H2SO4 10,030 6,930 6,720
Operating Cost, M $
Zinc Oxide Process 4,473 4,473 4,473
Plume Reheat 502 502 502
Sulfuric Acid 1. 745 1. 745 1.745
Total Operating Cost: 6,720 6, 720 6,720
Profit Before Tax, M$ 3,310 210 -0-
Profit After Tax, M$ 1,589 101 -0-
Return on Total Investment A/T, % 9.3 0.6 -0-
Payout years 5.0 9.0 9.5
Profit Before Tax
$/tonofcoal 0.93 0.06 -0-
mills/kwh 0.38 0.02 -0-
$/tonofH2SO4 n.22 1.61 -0-
Break-even Point 37% 66% 70%
-------
TABLE 105
ZINC OXIDE PROCESS: ECONOMIC ANALYSIS
Case 3
70% Plant Factor & 96% Conversion of SO2 to
I. Capital Investment
Zinc .Oxide Process
Plume Reheat
Sulfuric Acid
Working Capital
Total Investment:
$3,345, 700
71,800
1,354,000
272,300
—$5,044,000
II. Profitability
Sales, M$, 55. 4 M tons
Operating Cost, M $
Zinc Oxide Process
Plume Reheat
Sulfuric Acid
Total Operating Cost:
Profit (loss) Before Tax, M $
Profit After Tax, M $
Return on Total Investment A/T,'
Payout, years
Profit (loss) Before Tax
$/ton of coal
mills /kwh
$/ton of
Break-even Point
Sulfuric Acid, $/ton
$34 $23.50
1,884 1,302
1,949
1,949
(65)
-0-
-0-
(0.12)
(0.05)
(1.17)
(647)
-0-
-0-
(1.19)
(0.50)
(11.65)
73%
-------
TABLE 106
ZINC OXIDE PROCESS: ECONOMIC ANALYSIS
Case 4
90% Plant factor & 96% Conversion of SO, to SO.
I. Capital Investment
Zinc Oxide Process $8,784,400
Plume Reheat to 200°F 46, 100
Sulfuric Acid 3, 690, OOP
Sub-Total for Depreciation $12, 521, 500
Working Capital 655. 900
Total Investment «* $13,177,000
II. Profitability
Sulfuric Acid, $/ton
Sales. M$, 324 Mtons
Operating Cost, M$
Zinc Oxide Process
Plume Reheat
Sulfuric Acid
Total Operating Cost:
Profit Before Tax, M$
Profit After Tax, M$
Return on Total Investment, A/T,%
Payout, years
Profit, $/ton H2SO4> B/T
Break-even Point, %
-------
S2
JO
I
5
15
14
13
12
11
10
9
8
7
6
5
4
3
2
1
T—|—r
Expected level of
operations
emi-variable costs
Sales @$34/ton
Sales @ $24.85 /|
Sales @ 23.50/tor
90 100
10 20 30 40 50 60 70
Plant Factor - %,
168 252 ' 3^6 421
100% H2S04 - Thousand tons / Yr
ZINC OXIDE PROCESS - CASE 1
1400 MEGAWATT NEW POWER PLANT FACILITY
BREAK-EVEN CHART WITH CONVERSION OF S02 TO H2S04
Figure 48
-------
.2
I
to
o
Expected level of
Semi-variable costs _
Sales 0 $34 /ton
Sales @$23.50/ton
Sales® $22.78/ton
0 10 20 30 40 50 60 70 80 90 100
Plant Factor - %
0 84 168 252 336
100% H2S04 - Thousand tons/Yr
421
ZINC OXIDE PROCESS - CASE 2
1400 MEGAWATT EXISTING POWER PLANT FACILITY
BREAK-EVEN CHART WITH CONVERSION OF S02 TO H2S04
Figure 49
-------
JO
J.U
2.8
2.6
2.4
2.2
2.0
1.8
1.6
L2
1.0
.8
.6
.4
.2
0
-
•
Expected level of
operations
Break-
Vi
Sen
C
—
,
X
jriab
tots
^
i-vai
osts
—
/
y
-ever
e
N
^
Tabl
/
^
V
B^
/
7^
^
,^
~y
/
/
r
\\l
^
^
i
/
m ^MM
^
;edC
*
^.
^
/
y/
x
osts
\
J*
r
/
_ . .
/
*
/
/
—
_. _
/
r
^^
.
/+—
—
—
- Sales 6 $34/ton
Sales @ $23.50/ton
0 10 20 30 40 50 60 70 80 90 100
Plant Factor - %
0 16 32 48 64
100% H2S04 - Thousand Tons / Yr
79
ZINC OXIDE PROCESS-CASE 3
220 MEGAWATT EXISTING POWER PLANT FACILITY
BREAK-EVEN CHART WITH CONVERSION OF S02 TO H2S04
-------
15
14
13
12
11
10
9
sj, g
8
5
o
i i i
Expected level of
operations
Break-even
Fixed costs
I I I
Sales @$34/ton
7
6
5
4
3
2
1
0 10 20 30 40 50 60 70 80 90 100
Plant Factor, %
I | t | I I I I I I I
0 72 144 216 288 360
100% H2S04 - Thousand Tons/Yr
ZINC OXIDE PROCESS-CASE 4
NEW SMELTER FACILITY
BREAK-EVEN CHART WITH CONVERSION OF S02 TO H2S04
Figure 51
Sales @ $23.50/ton
Sales @$17.70/ton
-------
C. LIME PROCESS
1. Introduction
A reevaluation of the Mitsubishi Lime Process was made, as
reported earlier, which indicated that it was probably not economic to recover
gypsum as a salable by-product. For this reason it was decided that the lime
process would be treated as a simplified system which involved no product
recovery. This plan was not carried through to completion, however, since
other NAPCA investigators are giving adequate study to processes utilizing
lime and limestone. A Case 3 analysis for the Simplified Lime Process was
conducted, however, and it is this work which is summarized herein.
2. Process Design
A flooded bed absorber, similar to the Turbulent Contact
Absorber (TCA), was selected as the scrubbing system for this process. This
type of system consists of one or more beds in series of plastic spheres or
glass marbles in turbulent motion. The motion of the spheres and their
nonporous surface should prevent the accumulation of scale and thus eliminate
the need for a delay tank and expensive descaling operations as described in
the Howden-ICI Process.328
Although lime is more efficient, limestone was selected as
the absorbent on the basis of cost. An excess of 35% over the theoretical was
328 '
used to improve absorption. A three-stage unit was specified since field
tests indicated that only 80% of the SO- was removed from flue gas in a two-
343
stage scrubber us^ing limestone slurry as the absorbent.
It was assumed that this system would remove 90% of the SC>2
and 98-99% of the fly ash from flue gas. The slurry, consisting of CaCOj,
CaSO4- 2H2O, CaSOj- 2H2O, and fly ash is circulated at a high rate. A constant
make-up of CaCO^ is added with the simultaneous discharge of spent slurry to
-
Designed and manufactured by the Air Correction Division of Universal Oil
Products Company.
-------
a settling pond. The flow diagram, Figure 52, illustrates the system.
3. Capital Costs
The capital cost estimate, which is summarized in Table 107
indicated a total investment of $2, 022, 000, which is equivalent to a capital
requirement of $9. 19/kw capacity.
The purchased equipment cost was derived from the equipment
list (Appendix C) which shows basic specifications and estimated costs. The
derivation of working capital is given in Table 108.
4. Operating Costs
The operating cost at 70% plant factor was estimated at
$892, 000 per year, or $2. 08/ton of coal, and 0. 85 mill/kwh. The summary
is shown in Table 109.
The only raw material required is ground limestone at a rate
of 118,000 tons annually based on 98% purity and 0.2% in-plant losses.
Direct labor was based on one man per shift plus one man on
day shift for raw material handling.
The electrical load was estimated at 1790 kw. Make-up
water costs amount to $9, 400.
No cost was charged to waste disposal, since it was assumed
that adequate settling ponds would be available at the plant site.
5. Profitability
A profitability analysis was not applied to this system since
salable by-products are not generated.
V. RESULTS OF THE PHASE III EVALUATION
\
A. CAPITAL COSTS
Table 110 summarizes the total investment for each case. The capital
equipment is comprised of the specific process equipment, plume reheat system,
and sulfunc acid plant. The working capital is added to make up the total investment.
-------
PURIFIED GAS TO REHEAT OR STACK
LIMESTONE
<£>STREAM NUMBER
Q PRESSURE. INCHES HjO
[^TEMPERATURE, 'f
GAS
00
P-3
750 GPM
0.55 MM!
MMSCFM
BED
30.000 GAL
LIMESTONE
HOPPER / FEEDER
SOOLB/MIN
SLURRY MIXING
SO.OOO'GAL
ING
18
64
-
-
100
156
172
-
-
-
STREAM NUMBER
UNITS
HzO
SOz
TOTAL GAS
FLY ASH
CaCO)
CaSOj 2H20
CaSQj 2HjO
TOTAL SOLIDS
saurioN
GPM
1
moles
mln
95.9
3.97
1320
Ids
fflin
1730
254
31500
79.1
2
moles
•rin
1W
a 40
1390
Ibs
mln
3050
25.4
40600
3
moles
inin
17700
56
80.1
80.?
Ibs
min
319000
3500
5COO
1250D
13800
135400
319000
40000
4
moles
mln
389
L25
L78
1.78
Ibs
mln
7000
79.1
125
278
307
789
7000
891
5
moles
mln
311
Ibs
min
5600
5600
672
6
moles
min
4.82
Ibs
min
482
7
moles
min
311
4.82
Ibs
•rin
5600
482
482
5800
•72
8
motes
min
151
151
Ibs
min
2720
2720
327.
LOSSES
BS
min
LOO
LIME PROCESS FLOW DIAGRAM
-------
TABLE 107
LIME PROCESS: CAPITAL COST ESTIMATE SUMMARY
Case 3
ITEM COST - $
21. CAPITAL REQUIREMENTS: $ 9.19 /kw capacity
337
1. Purchased Equipment 468,000
2. Erection Labor 117,000
3. Foundations and Platforms 84. 200
4. Piping 234.000
5. Instruments 46, 800
6. Insulation 37, 400
7. Electrical 46, 800
8. Process Buildings, Structures 117, OOP
9. Plant Facilities, 5% of 1-8 57, 600
10. Plant Utilities, 7% of 1-9 84. 600
11. PHYSICAL PLANT COST . 1.293.400
12. Engineering & Construction, 20% of 11 258. 600
13. DIRECT PLANT COST 1.552.000
14. Contingency, 15% of 13 232.800
15. Contractor's Fee, 5% of 13 + 14 89.200
16. FIXED CAPITAL COST 1.874.000
17. Interest During Construction. 2. 5% of 16 ^ mm^TtQQQm
18. SUB -TOTAL FOR DEPRECIATION 1,921,000
19. Working Capital 101.000
-------
TABLE 108
LIME PROCESS: WORKING CAPITAL
Case 3
Plant Factor
COST - $
Raw Material Inventory, 2 months 28,800
Direct Labor, 3 months 8, 000
Maintenance, 3 months 14, 100
Supplies, 3 months 2,800
Payroll Burden, 3 months 1» 700
Plant Overhead, 4 months 17,400
Fixed Cost, 0. 5% fixed capital cost 9,400
Spare Parts & Miscellaneous, 1.0% fixed
capital cost
TOTAL
-------
TABLE 109
LIME PROCESS: OPERATING COST ESTIMATE SUMMARY
Case 3
70% Plant Factor
Fixed Capital Cost: $1,874,000
!
t
ITEM TOTAL $ %
1. Raw Materials fe Chemicals 381,000 42.71
2. Direct Labor 32. 100 3.60
3. Supervision, 15% of 2 4.900 0.55
4. Maintenance, 3% of fixed capital cost 56.200 6. 30
5. Supplies, 20% of 4 11.200 1.26
6. Utilities 78.700 8.82
7. Other
8. TOTAL DIRECT COST 564. 100 63.24
9. Payroll Burden, 18. 5% of 2 b 3 6.800 0. 76
10. Plant Overhead, 50% of 2, 3, 4 & 5 52.200 5.85
11. Waste Disposal - -
12. Other - -
13. TOTAL INDIRECT COST 5Q.QQQ 6.61
14. Depreciation, 11% depreciable capital cost 211. 300 23.69
15. Taxes & Insurance, 3% depreciable cap. cost 57.600 6.46
16. Other - -
17. TOTAL FIXED COST 268.900 30. 15
18. TOTAL OPERATING COST 892. OOP 100.0
19. Cost: $ 2.08 /ton coal, 0.85 mill/kwh
-------
TABLE
CAPITAL INVESTMENT SUMMARY
WITH CONVERSION OF SULFUR DIOXIDE TO SULFURIC ACID
Capital Investment, Thousand $
Process Equipment
Plume Reheat Equipment
Sulfunc Acid Plant
Total Plant
Working Capital
Total Investment
Capital Requirements, $/kw
Zinc Oxide Process
Case 1
15,085
205
4, 100
19,390
818
20,208
14.43
Case 2
11,925
205
4, 100
16,230
726
16,956
12. 11
Case 3
3,346
72
1,354
4,772
272
5, 044
22. 93
Case 4
8,785
46
3,690
12,521
656
13, 177
-
Simplified
Lime Process
Case 3
1,921
72
-
1,993
101
2, 094
-------
The capital requirement in terms of dollars per kilowatt of installed
capacity is the lowest for the Simplified Lime Process, Case 3. This process
is a relatively simple system since it does not have an absorbent regeneration
section, nor does it require a sulfuric acid plant for SO- recovery.
B. OPERATING COSTS AND PROFITABILITY
The initial economic analyses of the cases evaluated in Phase III
assumed a 90% plant factor (330 days operation per year). In Cases 1, 2 and 4
operations were considered to be continuous under these conditions. In Case 3,
however, operations averaged 60% of capacity on the operating days, equivalent
to a 54% plant factor. In all cases a plume reheat of 50°F (a rise from 122° to
172°F) was used.
The operating costs and profitability for these situations are summa-
rized in Table 111. It is interesting to note that the net cost of Case 3 (220 Mw)
of the Zinc Oxide Process is about the same as the cost of Case 3 of the Simplified
Lime Process, if acid from the ZnO Process were saleable at about $23/ton.
The changes in operating costs and profitability due to adjustments
in plant factor and plume reheat temperature are reflected in Table 112. The
data indicate that'operations in Cases 1 and 2 (1400 Mw) could break-even if
sulfuric acid could be sold for $20-25/ton. Case 3 operations for both the Zinc
Oxide and Simplified Lime Processes would operate at a substantial cost. Case
4, the New Smelter Facility, is not included in this tabulation since it would not
be affected by the 70% plant factor of power plants.
-------
TABLE 111
PROFITABILITY
PLANTS OPERATING AT 90% PLANT FACTOR - 330 DAYS PER YEAR
PLUME REHEAT FROM 122° TO 172°F
Oo
•£•
ro
Sales, MtonsH2SO4
Sales, M$
Operating Cost, M$
Process
Plume Reheat
Sulfuric Acid
Total
Profit (loss) Before Tax, M$
Profit After Tax, M$
Return on Total Inv A/T, %
Payout, years
Profit (loss) Before Tax
$/ton of coal
nulls/kwh
$/ton of H2SO4
Break-even Point, %
Zinc Oxide Process -
Case 1
34
380
12,920
5,645
422
1,925
7, 992
4,928
2,365
11 7
4 5
1 07
0 44
12.97
43
23.50
380
8,930
5,645
422
1,925
7.992
938
450
2 2
7.8
0 20
0 08
2.47
74
21.03
380
7.992
5,645
422
1,925
7,992
- 0 -
- 0 -
- 0 -
9.5
- 0 -
- 0 -
- 0 -
90
Case 2
Case 3*
•
Case 4
Sulfuric Acid Price, $/ton
34
380
12, 920
5,036
422
1,925
7,383
5,537
2,658
15 6
3 8
1.21
0 50
14.57
36
23 50
380
8,930
5,036
422
1,925
7.383
1,547
743
4 4
6.7
0 34
0.14
4.07
64
19.43
380
7.383
5.036
422
1.925
7.383
- 0 -
- 0 -
- 0 -
9-5
- 0 -
- 0 -
- 0 -
90
34
44
1.496
1, 137
90
652
1,879
(383)
-
-
-
(0 90)
(0. 37)
(0.87)
-
23 50
44
1,034
1.137
90
652
1.879
(845)
-
-
-
(1.98)
(0.81)
(1.92)
-
34
324
1.016
3,970
50
1,714
5,734
5.282
2,535
19.2
3.4
.
-
16.30
33
23.50
324
7.614
3.970
50
1.714
5,734
1,880
902
6.8
5.8
.
-
5.80
55
17.70
324
5,734
3,970
50
1,714
5,734
- 0 -
- 0 -
- 0 -
9.5
_
-
- 0 -
90
Lime
Case 3*
- 0 -
- 0 -
- 0 -
788
90
-
878
(878)
-
-
-
(2-05)
(0. 84)
-
-
Note. Also see Tables 26 and 28 of Volume U.
In Case 3, operations are at 330 days per year at 60% capacity.
**Assumes 95% SO2 removal efficiency, 10% conversion of removed S^ to a discardable sulfate stream,
-------
OO
^
UO
Sales, M tons
Sales, M$
Operating Cost, M $
Process
Plume Reheat
Sulfuric Acid
Total
Profit (loss) Before Tax,M$
-Profit After Tax, M$
Return on Total Inv , A/T, %
Payout , years
Profit (loss) Before Tax
$/ton of coal
mills /kwh
$/ton of
TABLE 112
PROFITABILITY
PLANTS OPERATING AT 70% PLANT FACTOR
PLUME REHEAT FROM 122° to 200°F
Break -even Point, %
Zinc Oxide Process
Case 1
Case 2
| Case 3
Lirritt
Case 3
Sulfuric Acid Price, $/ton
34
295
10,030
5,082
502
1,745
7,329
2,701
1,296
6.4
5.9
0 76
0.31
9 15
43
24 85
295
7.329
5,082
502
1,745
7,329
- 0 -
- 0 -
- 0 -
9.5
- 0 -
- 0 -
- 0 -
70
23 50
295
6,933
5,082
502
1,745
7,329
(396)
-
-
-
(0.11)
(0 05)
(0.17)
78
34
295
10,030
4,473
502
1,745
6,720
3,310
1,589
9 3
5.0
0 93
0.38
11.22
37
23 50
295
6,932
4,473
502
1,745
6,720
212
102
0.6
9.0
0 06
0.02
0.72
66
22.78
295
6,720
4,473
502
1,745
6,720
- 0 -
- 0 -
- 0 -
9.5
- 0 -
- 0 -
- 0 -
70
34
55 4
1,884
1,225
106
618
1,949
(65)
-
-
-
(0. 12)
(0. 05)
(1.17)
73
23 50
55 4
1,302
1,225
106
618
1.949
(647)
-
-
-
(1-17)
(0.48)
(11.68)
-
- 0 -
- 0 -
- 0 -
892
106
-
998
(998)
-
-
-
(1 81)
(0- 74)
-
-
Note
*
Also see Tables 25 and 27 of Volume IL
Assumes 95% SO^ removal efficiency, 10% conversion of removed SO- to a discardable sulfate stream,
-------
PART FIVE
"FUTURE" WORK
As an extension of the effort described in Parts Two to Four relating to
the removal of sulfur dioxide from flue gases by aqueous scrubbing methods,
the following specific areas of investigation were considered to warrant addi-
tional effort (and this work was conducted under Phase IV of this study; see
Volume II of this report for the results):
• Conceive new aqueous scrubbing processes (including the regeneration
step) for the removal of SO, from the flue gases emanating from
various industrial sources fof which, power plant and smelter effluents
will be considered representative). Through appropriate laboratory
investigations, develop these newly conceived processes by demon-
strating their technical feasibility.
• Through appropriate laboratory investigations, develop improvements
to previously-conceived aqueous scrubbing processes (or any portion
thereof). An example here could be an investigation of how to mimmizi
the effects of the disproportionation that occurs during the calcination
of metallic sulfites (a regeneration step applicable to various aqueous
processes).
• Through appropriate laboratory investigations, determine the degree
to which inadvertent sorbent oxidation in aqueous scrubbers can be
minimized by the utilization of various oxidation inhibitors and
complexing agents (both with and without fly ash being present in the
flue gas being tested).
Presuming that the degree of oxidation cannot be economically
reduced by use of inhibitors or complexing agents, investigate
the technical feasibility of separating the oxidation product
from the scrubber effluent by chemical or ion-exchange means,
'followed by the thermal, chemical or electrochemical regener-
ation of the resultant material.
Complete the study that was begun under the initial contract
that deals with the determination of the effects that pre-
scrubbing has on the degree of oxidation in the main SO-
scrubber.
• Through appropriate laboratory investigations, determine the degree
of interference which inadvertent sorption 6f NO into SOo scrubbing
solutions has on SO, removal efficiencies (both with and without fly
ash being present in the flue gas being tested).
-------
Assess the technical feasibility of achieving high NO removal
efficiencies in conjunction with high SO, removal efficiencies.
Support each of the above-mentioned laboratory investigations with
preliminary process evaluations/designs/economic analyses to
illustrate the economic feasibility of the concepts undergoing
scrutiny in the laboratory (using either the Phase I or the Phase III
approach of the initial contract, whichever is deemed appropriate).
-------
PART SIX
BIBLIOGRAPHY
During ttie course of the Phase I effort an extensive list of references
was acquired relating either directly or indirectly to some aspect of the aqueous
scrubbing of sulfur dioxide from flue gases. The bibliography which follows is
not intended to represent a complete coverage of the subject, but is believed to
include the most important papers. Of the first 645 entries approximately 500
(those marked by an asterisk) have been acquired and catalogued by Aerojet,
and many of these were utilized in the preparation of the present report.
The references have been categorized into six groups, which are con-
sidered in the following order: Articles, Patents (United States and Foreign),
Reports, Government Publications, Books, and Theses and Dissertations.
Within each group the listing is alphabetical by principal author. An index of
all subsidiary authors is also provided for completeness.
-------
ARTICLES
1. Adams, F.W., "The Absorption of SO, in Water, " Trans. Amer. Inst.
Chem. Eng., 28, 162.
2. Agliardi, N. and Slodyk, T. , "The Activated Adsorption of Sulfur Dioxide,
Oxygen, and Mixtures of the Two on Vanadium Oxide, " Gazz. chim. ital. ,
77, 66-75 (1947). CA:41-7197 (1947).
#
10(Z),20-2 (February 1968).
3.* Air Eng. , "New Developments in Industry for Pollution Control, "
(2) ,20-2 — '
4.* Albright, L. F. , Shannon, P. T., Yu, Sun-Nien and Chueh, Ping Lin,
"Solubility of Sulfur Dioxide in Polar Organic Solvents, " Chem. Eng.
Progr. Symp. Series, 59, 66-74 (1963).
*
5. Allen, L. N. , Jr. , "Recovery of Manganese from Low Grade Ores, "
Chem. Eng. Progr.. 50(1), 9-13 (January 1954).
6. Altybaev, M. and Streltsov, V. V. , "Removal of Sulfur Compounds from
Gaseous Fuels, " J. Air Poll. Control Assoc., Abstracts, No. 8355
(July 1967).
7. Amdur, M. O. , "Report on Tentative Ambient Air Standards for Sulfur
Dioxide and Sulfuric Acid, " Ann. Occupational Hyg., 3_, 71-83 (February
1961).
8. Andnanov, A. P. , "Combined Method of Purification of Flue Gases from
Oxides of Sulfur, " CA:31-3666(1937).
9. Andnanov, A. P. si Certkov, B.A., "Metoda Ciclica Amoniacala de
Captare a SO, din Gazele de Ardere, " Jurn. Prikl. Himii. 7_, 10 (July
1954). * -
JjC
10. Applebey, M. P. , "The Recovery of Sulfur from Smelter Gases, "
J. Soc. Chem. Ind. , _56_, 139T-46T (May 1937).
11.* Arai, K. , "Air Pollution Control in Petroleum Refineries, " Nenryo
Kyokaishi, 46(485), 669-79(September 1967)(in Japanese). CA:68-81179u.
12.* Arai, Y. , Takenouchi, H. and Nagai, S. , "Reactions Between Lime and
Sulfur Dioxide. I. Absorption Mechanism of SO? into Quick Lime,"
Sekko to Sekkai (Gypsum fa Lime), 4£, 5-11 (I960) (in Japanese, English
abstract).
13- Arnold, T. H. and Chilton, C. H. , "New Index Shows Plant Cost Trends, "
Chem. Eng. , 143-52(18 February 1963), or Chem. Eng. Report No. 224.
-------
14. Atsukawa, M. , "World Trend in SO, Removal Methods for Air Pollution
Control, " Kagaku Kogyo (Chemical fndustry) Tokyo, J.8J12), 11^5-92
(December 1967)(in Japanese).
15. Atsukawa, M. , Nishimoto, Y. and Matsumoto, K. , "Removal of SO2 Gas
from Waste Gases, " Mitsubishi Heavy Industries, Ltd. , - Tech. Rev..
2(2),51-7(May 1965).
16. Babushkina, M. D. , Babaev, E. V. and Molehanova, N.I., "Preparation of
Magnesium-Base Cooking Liquors, " Bumazhn. Prom. , 38(5), 7-10(1963)
(in Russian). CA:59-7746f.
17. Bachmair, A. , "Massnahmen zur Vehinderung von Luftverunreinigung
durch Dampfkraftwerke (Measures to Prevent Air Pollution by Steam
Power Plants). - Mitt. Vereinig, " Grosskesselbesitzer, 93.446-53(1964).
18.* Bahr, H. , "Das Katasulf-Verfahren, " Gluckauf. 73J40), 901-13(2 October
1937) (in German).
19. Baukloh, W and Valea, I. , "Effect of Sulfur Dioxide on Iron and Steel, "
Korrosion u. Metallsc'hutz, 15,295-8(1939).
20. Belcher, R. , "Dry Absorbents for the Absorption of Sulphur Dioxide,
with Special Reference to the Determination of Carbon in Steels, "
J. Soc. Chem. Ind. , 64, 111-4.
*
21. Bender, R. J. , "An Unusual Approach to Air Pollution Control, " Power,
njp_(12),83(December 1966).
22. Bender, R. J. , "Tall Stacks, A Potent Weapon in the Fight Against Air
Pollution, " Power, 111(12), 93-6(December 1967).
23. Benny, J.C. , "Sulphur Dioxide Recovery, " Pump and Paper Mag. Can.,
46_, 598. .
24. Bertolacini, R. J. and Barney, J. E. , "Colorimetric Determination of
Sulfate with Barium Chloramlate, " Anal. Chem. , 29, 281(1957).
25. Bettelheim, J. , Klimecek, R. , Strnad, M. and Chlumsky, F. , "Absorption
of Sulfur Dioxide in an Unpacked Column with Jets, " Chemicky Prumysl,
10_, 281-4(1960)(in Czech. , English abstract).
26. Beuschlein, W. L. , "The Recovery of SO-, " CA:31-2424(1937).
£•
27. Beuschlein, W. C. and Porter, M. A., "Sulphur Dioxide Over-All
Absorption," Paper Trade Journal, 111,43-6(19 December 1940).
-------
28. Bienstock, D. , Amsler, R. L. and Bauer, E.R. , Jr., "Formation
of Oxides of Nitrogen in Pulvenze-d Coal Combustion, " J. Air Poll.
Control Assoc. , 16(8), 442-5 (August 1966).
29. Bienstock, D. and Field, J. H., "Bench-Scale Investigation on
Removing Sulfur Dioxide from Flue Gases, " J. Air Poll. Control Assoc.,
10(2), 121-5 (April I960). -
30. Bienstock, D. , Field, J. H. and Benson, H.E., "Sulfur Dioxide in
Atmospheric Pollution, and Methods of Control: Chapter in Geo-
physical Monograph No. 3, Atmospheric Chem. of Chlorine and
Sulfur Compounds, " Am. Geophysical Union, Washington, D. C.,
54-62 (1959).
31. Bienstock, D. , Field, J. H. , Katell, S. and Plants, K. D. , "Evaluation
of Dry Processes for Removing Sulfur Dioxide from Power Plant Flue
Gases," J. Air Poll. Control Assoc. , 15(10), 459-64 (October 1965).
*
32. Bienstock, D. , Field, J. H. and Myers, J. G. , "Removal of Sulfur Oxides
from Flue Gas with Alkalized Alumina at Elevated Temperatures, "
J. Eng. Power, 86(3). 353-60 (July 1964).
33. Billinge, B. H. , "The Chemisorption of Sulfur Dioxide on Carbons, "
Conference, Industrial Carbon Graphite, Papers, 2nd, London, 398-404,
(1965).
34. Bonn, D. E. , "Wet-Type Dust Collectors, " Chem. Eng. Progr. ,
59(10), 69-74 (October 1963).
35. Bonsall, R. A., "Chart Gives Solubility of SO, in Ammonium Bisulfite, "
Chem. Eng. , 68(10), 182-4 (15 May 1961).
ifc
36. Bottoms, R. R. , "Organic Basis for Gas Purification. " Ind. Eng. Chem.,
23J5), 501-4 (May 1931).
37.* Boubhk, T. , Dvorak, K. , Hala, E. and Schauer, V- , "Gleichgewicht
Flussigkeit-Dampf in Systemen Elektrolytischer Komponentin IL
System NH, -SO,-SO,-H2O bei 50, 70, und 90°C, " Collection Czech.
Chem. Commun. , 28, 17^1-803 (1963) (in German).
38.* Bouvier, R. , "S"ifti* SiY»nfc« Removal System. " American Power Conf.
Proc. , 2£, 138-43 (1964).
39.* Brennan, P. J. , "Coal Researchers arc Grappling with Sulfur, " Chem.
Eng. , 74(21), 114-8 (9 October 1967).
40* Bretsznajder, S. , "Utilization of Sulfur Dioxide in Industrial Flue Gases,"
Prezemysl Chem. , 31(8), 276-84 (1952) (in Czech.). CA:47-10783d.
-------
41. Bretsznajder, S. , "Absorption of Sulfur Dioxide in Solutions of Basic
Aluminum Sulfate. V. Determination of the Extent of the Oxidation of,
the Absorbed Sulfur Dioxide by Oxygen, " Prezemysl Chem., 1.1, 115-7
(1955) (in Czech.). CA:53-831a.
42.* Bretsznajder, S. , Bacia, W. and Kawecki, W., "Absorption of Sulfur
Dioxide in Solutions of Basic Aluminum Sulfate. VIII. Absorption in
a Packed Column, " Prezemysl Chem., 35,83-6(1956) (in Czech.).
CA:51-17341f.
43.* Bretsznajder, S. , Kawecki, W., Kotowska, W. and Gantz, R., "Effect
of the Foam Layer on Bubble Absorption of Gases in Liquids, "
Prezemysl Chem., 35,564-5(1956) (in Czech.). CA:53-2702d.
A
44. Bretsznajder, S. , Kawecki, W. and Kotowska, W., "Mass Transfer
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-------
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$
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156. Galeano, S.F. and Harding, C.I. , "Sulfur Dioxide Removal and Recovery
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*
159. Ganz, S. N. , Kuznetsov, I.E. and Podgaiko, V. V. , "Sulfur Dioxide
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160. Ganz, S. N. , Kuznetsov, I.E., Shlifer, V.A. and Leykin, L. I. ,
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%
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178. Haagen-Smit, A. J. , "Studies of Air Pollution Control by Southern
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179. Haas, T. K. de. , Nieuwenhuizen, J. K. , Akbar, M. , Giessen, J.A.
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-------
180. Haley, H. E., "SO_ Removal Process Promises Cleaner Air, " Electrical
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*
181. Hammick, D. L. , "The Action of Sulfur Dioxide on Metal Oxides, "
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*
182. Hangebrauck, R. P. and Spaite, P. W. , "Controlling the Oxides of Sulfur, "
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*
183. Haemsch, E.and Schroeder, M. , "The Recovery of Sulphurous Acid From
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jj-
184. Hargrave, J. H. D. and Snowball, A. F. , "Recovery of Fume and Dust
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185. Harris, D. N. , "Reducing Sulfur Emissions," Combustion, 39(5), 36-8
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JJC
186. Haselbarth, J. E. and Berk, J. M. , "No. 31: Chemical Plant Cost
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187. Heitmann, H. G. and Sieth, J. , "Entschwefelung von Rauchgasen, "
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188. Heller, A. N. and Walter, D. F. , "Impact of Changing Patterns of
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189. Henrich, J. , "Practical Experiences in Removing SO, From Effluents
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190. Herzog, G. , "Desulfunzation of Flue Gases, Problems and Solutions,"
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192.* Hewson, G.W., Pearce, S. L. , Pollitt, A. and Rees, R. L. , "The
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-------
193. Higashi, M., Fukui, S. and Kamei, K. , "Study and Experience ef
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194. Hitchcock, L. B. and Scribner, A.K. , "Anhydroud Liquid SQ2," ted. Eftfl.
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195. Hofman, H. O. and Wanjukow, W., "The Decomposition of Metallic
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196.* Holzl, F., "The Ternary System K?O-SO,-H,O, " Z. ElektrQchem,
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197. Howat, D. D., "Removal and Recovery of Sulphur From Smelter Qageti'1
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198.* Huff, W. J. and Logan, L., "The Purification of Commercial Gases at
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200. Ingraham, T. R. and Kellogg, H. H., "Thermodynamic Properties of
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*
201. Ivanov, D. and Kostadinov, N., "Simultaneous Absorption of Hydrogen
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_H(3), 53-6'4 (1964) (in Russian). £A:66-443p.
202.* Jackson, A. and Solbett, J. M., "Sulphuric Acid Plant: Tail Gas
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203. Jarosz, A. , "Desulfurization of the Gas by the Wet Thylox Method, "
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204. Jellinek, K. , "The Electrolytic Preparation of Hyposulfite From
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-------
205. Jenness, L. C. and Caulfield, J. G. L., "Absorption of SO- in Water;
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102,37-41 (28 December 1939). —
206. Johnson, J. E. , "Gas Cleaning with Scrubbers, " J. Metals, 17(6), 670-2
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*
207. Johnson, J. E. .- "Wet Washing of Open Hearth Gases, " Iron Steel Eng. ,
44,96-8 (February 1967).
208. Johnstone, H. F. , "Reactions of Sulfur Compounds in Boiler Furnaces, "
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209. Johnstone, H. F. , "Metallic Ions as Catalysts for the Removal of Sulfur
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(May 1931). —
*
210. Johnstone, H. F. , "Progress in the Removal of Sulfur Compounds From
Waste Gases," Combustion, _5, 19-30 (August 1933).
#
211. Johnstone, H. F. , "Recovery of Sulfur Dioxide From Waste Gases -
Equilibrium Partial Vapor Pressures Over Solutions of the Ammonia-
Sulfur Dioxide-Water System, " Ind. Eng. Chem. , 27(5), 587-93
(May 1935).
212.* Johnstone, H. F. , "Recovery of Sulfur Dioxide From Waste Gases -
Effect of Solvent Concentration on Capacity and Steam Requirements
of Ammonium Sulfite-Bisulfite Solutions, " Ind. Eng. Chem. ,
2JH12), 1396-8 (December 1937).
213.* Johnstone, H. F. , "Recovery of Sulfur Dioxide From Dilute Gases, "
Pulp and Paper Mag. Can. , 53, 105-12 (March 1952).
214. Johnstone, H. F. , "Properties and Behavior of Air Contaminants,"
Proc. of U.S. Tech. Conf. on Air Pollution. 156-66, McGraw-Hill,
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#
215. Johnstone, H. F. and Coughanowr, D. R. , "Absorption of Sulfur
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%
216. Johnstone, H. F. and Keyes, D. B. , "Recovery of Sulfur Dioxide From
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217. Johnstone, H. F. and Klemschmidt, R.V., "The Absorption of Gases
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-------
218. John a tone, H. F. and Leppla, P. W., "The Solubility of Sulfur Dioxide
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219. Johnstone, H. F. and Moll, A.J. , "Air Pollution: Formation of
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220. Johnstone,' H. F. and Singh, A. D. , "Recovery of Sulfur Dioxide From
Waste Gases: Design of Scrubbers for Large Quantities of Gases,"
Ind. Eng. Chem. , £9_(3), 286-97 (March 1937).
221. Johnstone,' H. F. and Singh, A. D. , "Recovery of SO, From Waste
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Ind. Eng. Chem. , 32_, 1037-49 (1940).
222. Johnstone, H. F. and West, W.E. , "Recovery of Sulfur Dioxide From
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223. Johnstone, H. F. and Winsche, W.F., "Fused Salt Mixtures as Reaction
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224. Johnstone, H. F. , Read, H. J. and Blankmeyer, H. C. , "Recovery of
Sulfur Dioxide From Waste Gases: Equilibrium Vapor Pressures Over
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*
225. Johswich, F. , "The Desulfurization of Waste Gases - Importance and
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Steam Boilers by the Clean Air Process, " VIK Reports, 155, 2-19
(August 1964).
226.* Johswich, F., "The Present Position of Flue-Gas Desulfurization,"
Brennstoff - Waerme-Kraft. 1J(5), 238-45 (1965); Combustion
(October 1 $65).
227. Junge, C.E., "Sulfur in the Atmosphere, " J. Geophys. Res., 65,227
(1960). """
228. Junge, C.E. and Ryan, T. G. , "Study of the SO, Oxidation in Solution
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84_,46 (i958).
229.* Juntgen, H. and Peters, W., "Technical Principles of Separating
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230. Juntgen, H. , "Flue Gas Desulfurization, " Staub (Reinhaltung Luft),
28_(3), 89-93 (1968) (in German). CA:68 - 45838y.
-------
231. Kalushm, A.E. , Leont'eva, L..S. and Kas'yan, D. T. , "Absorption
of Sulfur Dioxide Under Froth Conditions (I), " Tr. Vses. Neft Nauchn.
Issled. Inst. po Tekhn. Bezopasnosit. 16, 109-13 (1964) (In Russian)
Abstract Only. CA:64 - 926W. —
#
232. Katell, S. , "Removing Sulfur Dioxide From Flue Gases, " Chem. Eng.
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233. Katell, S. and Plants, K. D., "Here's What SO. Removal Costs, "
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*
234. Katz, M. and Cole, R. J. , "Recovery of Sulfur Compounds From
Atmospheric Contaminants, " Ind. Eng. Chem. , 42(11), 2258-69
(November 1950). —
235. Kawazoe, K. , "Removal of Sulfur Dioxide From Flue Gases, "
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jjt
236. Kay K. , "Air Pollution Review 1956-57," Ind. Eng. Chem. ,
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237. Kennaway, T. , "The Fulham-Simon-Carves Process for the Recovery
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238. Kennaway, T. , Wood, C.W. and Box, P. L. , "A New Development
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239. Ketov, A. N. and Shhgerskii, A.S., "Laboratory Testing Methods -
Dry Lime Method for Removing Sulfur Dioxide From Heat and Electric
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240. Kettner, H. , "Removal of Sulfur Dioxide From Effluent Gases, "
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241. Kettner: H. , "The Removal of Sulfur Dioxide From Flue Gases, "
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242* Khokhlov, S.F., Annenkov, V.A. and Shutkin, G.A., "Mass Transfer
in a Scrubber With Cone-Shaped Grid Plates, " Khim. i Neft. Mashinostr,
9,25-6 (1965) (in Russian). CA:64-3075g.
243.* Kielback, A.W- and Crampton, E.W., "Progress by the Aluminum
Company of Canada, Limited, in Air Pollution Control, " presented
at the National Conference on Pollution and Our Environment. Montreal,
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-------
244. Kim, M. Rin and Bang, Oo Hoon, "Mechanism of Reaction Between
Manganese Oxides and Sulfur Dioxide in Aqueous Solution, " Chosun
Kwahakwo Tongbo, 4,43-8 (1964) (in Korean). Abstract Only.
CA:bZ-15745f.
245. King, R. A. , "Ecojiomic Utilization of Sulfur Dioxide From Metallurgical
Gases, " Ind. Eng. Chem. , 42J11), 2241-8 (November 1950).
246. Kirkpatrick, S. D. , "Trail Solves its Sulphur Problem, " Chem. Met. Eng.
45_, 483-5 (September 1938).
247. , Kishinevsky, M. Kh. , and Fayer, S. M. , "Kinetics of Absorption of
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248. Kiyoura, R. , "Studies on the Removal of Sulfur Dioxide From Hot Flue
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^6_(9), 488 -9 (September 1966).
249. Kiyoura, R. , Kironuma, H. and Uwanishi, G. , "The Recovery of Sulfur
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Inst. Technol. . 8^,1-5(1967) (in English). CA:68 - 98412n.
250.* Kleinschmidt, R. V- , "Flue Gases Laundered to Prevent Air Pollution,"
Power Plant Eng. , 42_, 393-6 (June 1938)
251. Klimecekj R. , "Czechoslovakia!! Proposal of Ammoniacal Flue Gas
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3_, 175-9 (1967) (in English). CA:68 - 69527w.
252. Klimecek, R. , u. Bettelheim, J. , "Absorptionskolonne mit Schauben-
formiger Drahtfullung, " Zh. Prikl. Khim. , 36^, 2432-7 (1963).
253. Klimecek, R. , Skrivanek, J. and Bettelheim, J. , "Desulfuring Flue
Gas," Staub. 26(6), 235-8 (1966) (in German).
*
254. Knott, K. H. and Tuerkoelmez, S. , "Krupp Rotary-Brush Scrubber
for Control of Gas, Vapor, Mist, and Dust Emissions, " Tech. Mitt.
Krupp. Werksber, 24_(1),25~8 (1956) (in German). CA:65 - 10149g.
255. Kohl, A.L. and Riesenfeld, F.C., "Gas Purification: Sulfur Dioxide
Removal by Liquid Absorption, " Chem. Eng. , 66(12), 147-51 (15 June
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256. Kohler, K. H. ."Moglichkeiten eines Betriebsvergleichs von Gasentschwe-
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-------
257. Kopita, R. and Gleason, T.G., "Wet Scrubbing of Boiler Flue Gas, "
Chem. Eng. Proqr. , 64(1), 74-8 (January 1968).
258. Krishman, V. S. R. , "Removal of Sulfur Dioxide From Stack Gases of
Sulfuric Acid Plants," Technology (Sindri), Spec. Issue, 3(4), 51-3
(1966) (in English), CA:68 -89701d."~
259. Kriz, M. , "Desulfunzation of Flue Gases by Calcium Carbonate at
High Temperatures, " Pr. Ustavu Vyzk, Paliv. , 15, 106-28 (1967)
(in Czech. )• CA:68 -45839w.
*
260. Kronseder, J. G. , "Economics of Phosphoric Acid Processes, "
Chem. Eng. Progr. , 64(9), 99 (September 1968).
*
261. Kropp, E.P. and Simonsen, R. N. , "Scrubbing Devices for Air Pollution
Control, " Air Poll. Smoke Prey. Assoc. Proc. , 45,48-53 (1952).
262. Kruel, M. and Juntgen, H. , "On the Reaction of Calcined Dolomite and
Other Alkaline Earth Compounds With the Sulfur Dioxide of Combustion
Gases as Carried out in a Cloud of Suspended Dust,!l Chemie Ing. Tech.,
39_, 607-13 (1967).
263. Kulcsar, G. J. and Lengyel-Szabo, G , "The System Sulfur Dioxide Aniline.
III. Absorption Isotherms in Aqueous Solution, '' Studia Univ. Babes-Bolyai,
Ser. Chemia, 9_(1), 77-83(1964) Abstract Only. CA:61-15403a.
264.* Kuzmmykh, I. N. , Popov, D. M. and Gorbachev, B. I. , "Absorption
of Sulfur Dioxide in Perforated Plate Towers to Obtain a Strong
Ammonium Bisulfite Solution, " Khim. Prom. , 2, 128-32, 160
(in Russian). CA:56 - 7087a.
265. Kuznetsov, I.E. and Ganz, S. N. , "Purification of Industrial Gases From
Sulfur Dioxide, " Izv. Vysshikh Uchebn. Zavedenii, Khim. i Khim. Tekhnol. ,
9(1), 89-93 (1966). CA:65 - 8379*.
3&C
266. Laberge, J. C. , "Sulfite-MgO System -Sulfur Dioxide Absorption
Efficiency Improvement, " Tappi. 46(9), 538-41 (September 1963).
*
267. Laffey, W. T. and Manning, R. N. , "Solvent Selection for the
Reduction of Air Pollution, " Hercules Chemist. 56, 1-6 (March 1968).
268.* Lang, H. J. , "Simplified Approach to Preliminary Cost Estimates, "
Chem. Eng. (June 1948).
-------
269. Lapple, C.E. and Kamack, H. J., "Performance of Wet DustAScrubbers,"
Chem. Eng. Progr., 51J3), 110-21 (1955). *
270. Laubusch, E. J., "Sulfur Dioxide: Properties, Methods of Handling and
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Public Works, 9£, 117-21 (August 1963).
271. Lawler, C., "Air Pollution Control by Sulfur Dioxide Scrubbing System,"
J. Air Poll. Control Assoc., 1(1), 29-30 (May 1957).
272. Leclerc, £., "The State in Which Sulfur Exists in the Combustion Gases
From Furnaces, " ISeme Congr. Chim. Ind., Brussels, 858-69 (September
1935). CA:30-5756. ' l
273. Lepsoe, R., "Chemistry of Sulfur Dioxide Reduction - Kinetics and
Thermodynamics, " Ind. Eng. Chem., 3£, 92-100 (1938).
274. Lepsoe, R. and Kirkpatrick, W.S. , "Sulphur Dioxide Recovery at Trail,"
Trans. Canadian Inst. Min. and Met., 40, 399-404 (1937).
275. Lepsoe, R. and Kirkpatrick, W.S. , "Recovery of Sulphur From Sulphur
Dioxide, " Consolidated Mining and Smelting Company of Canada.
Pulp and Paper Mag. Can., 39, 20-1 (January "1938).
276. Leasing, R. , "The Development of a Process of Flue Gas Washing
Without Effluent, " J. Soc. Chem. Ind., 5£, 373-88 (1938).
277. Leasing, R., "Elimination of Sulphur From Flue Gases, " Engineering,
^46,499-501 (1938).
278. Lichtenstein, S., "Inside Air Pollution, " Mech. Eng., 89(11), 61-3
(November 1967).
279. Lisle, E. S. and Sensenbaugh, J. D. , "The Determination of Sulfur
Trioxide and Acid Dew Point in Flue Gases, " Combustion, 36, 12-6
(January 1965).
280. Lucas, D. H. , Moore, D. J. and Spurr, G., "The Rise of Hot Plumes
From Chimneys, " Int. J. Air Water Poll. , 7,473-500 (August 1963).
*
281. Ludwig, J. H. , "The Future in Air Pollution Control, " Heating, Piping,
Air Cond. , 39J12),65-8 (December 1967).
*
282. Ludwig, J. H. and Spaite, P. W., "Control of Sulfur Oxide Pollution, "
Chem. Eng. Progr. , 63_(6) (June 1967).
*
283. Ludwig, J.H. and Steigerwald, B. J. , "Research in Air Pollution:
Current Trends, " Amer. J. Public Health, 55(7), 1082-92 (July 1965).
-------
284. Ludwig, S., "Antipollution Process Uses Absorbent to Remove SO
From Flue Gases, " Chem. Eng., 74(3), 70-2 (29 January 1968). 2
(Process flowsheet).
*
285. Lukacs, J. and Rossano, A. T. , "Air Pollution and its Control, "
J. Can. Petrol. Technol. , Montreal, _6(l),23-6 (January-March 1967).
286. Mader, P.P. Hamming, W. J. and Bellin A. , "Determination of Small
Amounts of Sulfunc Acid in the Atmosphere, " Anal. Chem. , 22, 1161
(1950). ~
*
287. Maksimov, V. F. , Bushmelev, V.A. and Isaeva, N. M. , "Sorption of
H2S and SO2 by a Weak White Liquor Under Turbulent Conditions, "
Tr. Leningu.Tekhn6l. Inst. Tsellyulozn, - Bumazhn. Prom. ,
1£, 80-5 (1965) (in Russian). Abstract Only. CA:65 -3377d.
*
288. Mallatt, R.C., "Product Sulfur Reductions -- Expenditures and Results
Petroleum Industry, " presented at 60th Annual Meeting of the Air
Pollution Control Association (11-16 June 1967).
289. Manvelyan, M. G. , Grigoryan, G. O. and Gazaryan, S.A. , "Adsorption
of Mixtures of SO, and N Oxides by CaCO, Suspensions, " CA:54-9226c
(I960). * *
290. Manvelyan, M. G. , Grigoryan, G. O. , Gazaryan, S.A., Papyan, G. S. ,
Grigoryan, M. M. and Mirumyan, R. L. , "Sulfur Removal in Gasifying
Solid Fuel by Ca(OH)?; SO2 Removal From N2 Oxide-Containing Flue
Gases by Absorption in Mg(OH)2> " CA:55-11775c (1961).
291. Marchal, G. , "Thermal Decomposition of Sulfates, " Jour.chim. phys. ,
Z2, 559-82 (1925).
292* Markant, H. P. , Phillips, N. D. and Shah, I. S. , "Physical and
Chemical Properties of Magnesia - Base Pulping Solutions, "
Tappi, j!8(ll), 648-53 (November 1965).
293. Martin, A. and Barber, F. R. , "Investigations of Sulphur Dioxide
Pollution Around a Modern Power Station, " J. Inst. Fuel, 39,294-307
(July 1966).
294. McCabe, L. C. , "Significance of Sulfur Dioxide as an Air Contaminant, "
Proc. Amer. Power Conf. , 18,201-5(1956).
-------
295. McGavack, J. and Patrick, W.A. , "The Adsorption of Sulfur Dioxide
by the Gel of Silicic Acid, " J. Am. Chem. Soc. , 42, 946-78 (1920).
296. McLaughlin, Jr. , J. F., "Atmospheric Pollution Considerations
Affecting the Ultimate Capacity of a Thermal Electric Power Plant
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297. McPhee, D. T., "Powe.r Plant Using High Sulfur Coal Takes Steps to
Reduce Air Pollution, " Ind. Water Wastes, 8_, 9-11 (January-February
1963).
298. Meethan, A. R. , "Natural Removal of Pollution From the Atmosphere, "
Quart. J. Royal Met. Soc.. 76,359(1950).
299. Michaelis, P. , "Review of German Papers From 1926 to 1941 on Gas
Purification Methods - Alkazid, Thylox, Polythionate, Katasadf.,
Pel u Kohle. 37. 949-55 (1941). CA:36-3025.
300. Miller, D. M. and JonaHn, J. , "Kansas P & L to Trap Sulfur with
Flue-Gas Scrubber, " Electrical World, (4 March 1968).
301. Mirev, D., Elenkow, D. and Balarev, K., "Effect of Surface Active
Agents on the Rate of Absorption, " Izv.inst. obshcha. neorg. khim.
org. khim., Bui gar ahad. nauk,' 8, 73 -81 (1961) (in Rus sian).
302. Misaka, Y. and Yamade, N., "Absorption of Sulfur Dioxide by Calcium
Hydroxide Slurry, " Kagaku Kogaku. 31(9), 925-6 (1967) (in Japanese).
CA:68, 32953h. —
#
303. Monkhouse, A.C. and Newall, H. E. , "Industrial Gases - Recovery of
Sulfur Dioxide, " presented at Conference at Sheffield University
Disposal of Industrial Waste Materials, 103-7 (17-19 April 1956),
304. Moses, H., Carson, J. E. and Strom, G. H., "Effects of Meteorological
and Engineering Factors on Stack Plume Rise, " Nuclear Safety,
6(1), 1-19 (1964).
#
305. Nakagawa, Shikazo, "Removal and Utilization of Sulfur Dioxide in
Stack Gas by the JECCO Process, " Ryusan, 16_, 211-8 (1963) (in
Japanese). CA:60-10251c.
306. Napier, D. H. and Stone, M. H. , "Catalytic Oxidation of Sulfur Dioxide
at Low Concentrations, " J. Applied Chem. , 8, 781-6 (December 1958).
-------
307. Nelson, H. W. and Lyons, C.J. , "Sources and Control of Sulfur -
Bearing Pollutant, " J. Air Poll. Control Assoc. , 7(3), 187-93
(November 1957). ~~~ -
308. Newall, H. E. , "The Ammonia Process for the Removal of Sulfur
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309. Newall, H. E. and Eaves, A. , "The Effect of Wind Speed and Rainfall
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*
310. Nikolaev, A.M., Safin, R. Sh. and Karasev, A. G. , "Mass Transfer
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*
311. Nilsen, J. , "Air Pollution: Costly to Ignore, Costly to Control, "
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*
312. Nonhebel, G. , "A Commercial Plant for the Removal of Smoke and
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-------
319. Parker, A., "Atmospheric Pollution; Cost of Flue Gas Washing, "
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320.* Parker, C.H., "Plastics and Air Pollution. " SPE Journal. 23J12), 26-30
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321. Parkison, R. V. , "The Solubility of Sulphur Dioxide in Water and
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*
322. Parkison, R.V-, "The Absorption of Sulphur Dioxide From Gases of
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*
323. Paul, B. B. and Mitra, A.K. , "Recovery of Sulphur Dioxide From
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•
324.* Pawlikowski, S. , Aniol, S. and Bistron, S. , "Problems of Ammonia -
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*
325. Pawlikowski, S. , Szaraware, J. and Synoradzki, Z., "Stabilization of
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#
327. Pearson, D. A., Lundberg, L.A., West, F. B. and McCarthy, J. L.,
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*
328. Pearson, J. L. , Nonhebel, G. and Ulander, P. H. N. , "The Removal of
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329. Pearson, J. L. , Nonhebel, G. and Ulander, P. H. N. , "The Removal of
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*
330. Pechkovsku, B. B. and Kenroe, W.H. , "Investigation of the Thermal
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331. Pechkovsku, B. B. and Ketov, A. N. , "Study of the Thermal Decompo-
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-------
332. Peisahov & Chertkov, B. A. , "Report on Work Done by Several USSR
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333. Perry, H. , "Oxides of Sulfur and the Electric Utility Industry, "
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jjj
334. Perry, H. and Field, J. H. , "Air Pollution and the Coal Industry, "
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*
336. Phillips, C.W. and Dickey, S.W., "Air Pollution Control Features of
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*
337. Pmaev, V. A. , "Method for Continuous Extraction of Slime From a
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Vest. Techn. i Ekon. Inform. Nauchn-Issled. Inst. Techn.-Ekon.
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*
338. Pinaev, V. A. , "Partial Pressure of SO, Over Solutions of Magnesium
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*
339. Pinaev, V.A. . Pitelma, N. P. , Novikov, A.I. and Sosekina, G. V. ,
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340.* Plekhotkin, V.F., Kitts, A. P. and Gavlovskaya, S.S., "Elimination
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341* Plumley, A. L. , Jonakin, J. , Whiddon, O. D. and Shutko, F.W..
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343.* Pollock, W.A., Tomany, J. P. and Fneling, G. , "Flue Gas Scrubber, "
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-------
344. Pollock, W.A., Tomany, J. P., Frieling, G., "Sulfur Dioxide and
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345* Potop, P., "Reclaiming Sulfur Dioxide From Waste Gases, " Rev.
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346 * Potop, P., Creanga, L. and Teodorescu, C. , "Recovery of Sulfur
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348. Potter, A.E. , Harrington, R. E. and Spaite, P. W., "Limestone-
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350. Pozin, M. E., "Absorption of Sulphur Dioxide by a Sodium Carbonate
(Na_C
-------
*
355. Reed, L. E., "The Removal of Sulfur Dioxide From Flue Gas, "
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*
356. Rees, R. L. ,"The Removal of Oxides of Sulfur From Flue Gases, "
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*
358. Rees, R. L. , "Present Performance and Scope for Improvement in
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360. Remy, H. and Hene, W- , "The Adsorption of Gases by Active Charcoal, "
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361. Renzetti, N. A. and Doyle, G. J. , "Photochemical Aerosol Formation
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364. Riou, P. and Berard, P. A. , "The Rate of Absorption of SO2 by Alkaline
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365. Robins, D. L. and Mattia, M. M. , "Computer Program Helps Design
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366. Roesner G. , "The Sulfidine Method, A New Means of Utilization of
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367. Rohrman, F.A. and Ludwig, J. H. , "Sources of Sulfur Dioxide Pollution,"
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368* Rohrman, F.A., Steigerwald, B. J. and Ludwig, J. H. , "SO., Pollution:
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-------
369. Rosenbled, C., "Recovery of Heat and Sulphur Dioxide Gas in Sulphite
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*
370. Ross, C.R. and Rispler, L. , "Air Pollution Control in Canada,"
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*
371. Ross, L. W. and Lewis, H. C., "The Reaction of Sulfur Oxides with
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*
372. Rossano, A. T. and JLukacs, J. , "Air Pollution and its Control,"
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373.* Rueb, F. , "Procedures and Installations for Neutralizing Toxic
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* '
374. Russell, E. J. and Smith, N. , "The Combination of Sulphur Dioxide
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375.* Ryason, P. R. and Harkins, J. , "A Method of Removing Potentially
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#
376. Ryason, P. R. and Harkins, J. , "Studies on a New Method of Simul-
taneously Removing Sulfur Dioxide and Oxides of Nitrogen From
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377." Saenz, O. , McKee, H. C. , Hiser, L. L., and Reinauer, T.V.,
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378. Safin, R. Sh. , Zhavoronkov, N. M. and Nikolaev, A.M., "Study of
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379. Scheidel, C. , "Zur Beseitigung Anorganischer Emissionen in der
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380. Schnell, H. , "The New Plant for Production of Sulfur Dioxide at
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-------
381. Schwarz, K. , "Sulfur Dioxide Emissions, " Staub, 21, 71-7 (1 February
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*
382. Scott, W. and McCarthy, J. L. , "The System Sulfur Dioxide -
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383. Segal, A. Ya. , "Cleaning Waste Gases in the Dorogomilovsk Chemical
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384. Seidman, E. B. , "Determination of Sulfur Oxides in Stack Gases, "
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385. Sensenbaugh, J. D. , "Formation and Control of Sulfur Oxides in Boilers, "
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#
386. Shah, I.S. , "New Flue-Gas Scrubbing System Reduces Air Pollution, "
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387. Sherwood, T. K. , "Solubilities of Sulfur Dioxide and Ammonia in Water, "
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jjC
388. Simo, J. B. and Novella, E.G., "Mass Transfer in Absorption Processes,"
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*
389. Sknvanek, J. and Cada, V. , "Absorption of Sulphurous Oxide in a Venturi
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*
390. Slack, A. V. , "Air Pollution: The Control of SO2 From Power Stacks.
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391. Smalley, G. E. and Klohr, J. W., "Refinery Sulfur Recovery Aids Air
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392. Smalley, G. E. and Klohr, J. W., "Recovery of Refinery Sulfur in Air
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%
393. Smith, J. L. and Snell, H.A., "Selecting Dust Collectors, " Chem. Eng.
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394.* Smith, M. E. , "The Tall Stack, " Mech. Eng. , 90_(2),20-2 (February 1968).
-------
395. Snow. R. D., "Conversion of Coal Sulphur to Volatile Sulphur Compound
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399.* Spalding, C. W- and Han, S. T. , "Absorption With Chemical Reaction
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402. Spitalnjk, Z. and Stary, M. , "The Absorption of SO, From Flue
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403. Squires, A. M. , "Cyclic Use of Calcined Dolomite to Desulfurize
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404.* Squires, A. M. , "Air Pollution: The Control of SO, from Power
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405.* Squires, A. M. , "Air Pollution: The Control of SO, From Power
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jjc
406. Squires, A. M. , "Air Pollution: The Control of SO, From Power
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407. Stalker, W.W., Dickerson, R.C. and Kramer, G. D. , "Atmospheric
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-------
408* Stavivo, "Calcium Sulfate Produced by the Sulfurization of Thermal
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409. Stern, A. C. , "Survey of Air Pollution Research in Europe, " Amer.
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jje
410. Stites, G. , Jr. , "The Catalytic Oxidation Process for Removing
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411. Stollery, J. L. , "Fundamentals of Fluid Bed Roasting, " Eng. MiningJ. ,
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412. Stone, G. N. and Clarke, A. J. , "British Experience with Tall Stacks
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413. Stourdze-Visconti, Y. , "Obtaining Elementary Sulfur. The Catalytic
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415. Sutton, O. G. , "The Theoretical Distribution of Airborne Pollution
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*
417. Sutton, P. , "Air Pollution in Petroleum Refining, " Pt.U, Chem. Process
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*
419. Szarawara, J. , "Studies on the Statics of a System Water-Ammonia -
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Sj{
42°- Tailor, J. P. , "The Dual Cycle Regenerative Process for the SO-
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-------
7f
421. Takeucty, N., "Study of the Absorption of Sulfur Dioxide by a Liquid
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#
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431*
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427. Taylor, L. F., "Neutralization of SO," Ind. Chemist, 16,325-8
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£
428. Terraglio, F. P. and Manganelli, R. M. , "The Absorption of Atmos-
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432. Tieman, J.W. and Cylmer, A. B. , "The Gas-Phase Oxidation of Sulfur
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433. Tomkinson, M. G. , "Catalytic Hydrogenation of Sulfur Dioxide, "
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434. Torres, R. and Hahn, E. , "Studies of the Burkheiser Ammonium
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-------
435. Trautz, M. and Helfrich, F. , "Influence of Impurities and Applicability
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%
436. Trendafelov, D. , Ponyankov, B. and Zapryanova, A. , "Extraction of
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437. Tsudo, K. , "Operation of SO, Gas Absorbing Plant, " Ryusan,
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438. Uno, T. , Yamada, H. , Higashi, M. Fukui, S. and Atsukawa, M. ,
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439.* Vedensky, D. N. , "How the SO- Process Worked on Three Kids
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443.* Wahnschaffe, E. , "Desulfurization of Smoke Gases by the Clear Air
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-------
446.* Wallis, E., "Recovery of Sulphur in Marketable Form From Flue
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447. Wallsom, H. E., "Elimination of Dust and Sulphur Fr6m Boiler Flue
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*
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-------
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*
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#
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$
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-------
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#
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-------
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*
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*
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-------
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1
-------
#
548. Newall, H. E. , "Removal of Oxides of Sulphur From Flue Gases;
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#
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4
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-------
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-------
*
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*
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-------
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#
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Proctor and Gamble Company (Internal report), 1968.
#
602. Plumley, A. L. , Jonakin, J. , Martin, J. R. and Singer, J. G. , Removal
of SO, and Dust From Stack Gases., A Progress Report on the C-E Air
Pollution Control System, Southeastern Electric Exchange, 11-12 April
1968.
603. Power, Air Pollution; Special Report, August 1965.
604. Reid, W. T. , Recommendation for Use of Limestone and Dolomite in
Boiler Furnaces, Battelle Memorial Institute Research Rept. . Contract
PH 86-66-108, 30 June 1966.
*
605. Stanford Research Institute, Air Pollution Control, Report No. 353,
August 1968.
606. Stemkohlen Elekrizitat A. G. , Interim Report on the Status of Development
Work in the Field of Sulfur Removal From Flue Gases by the Additive
Process. .December 1965 and June 1966. ~~~~
607. Tanaka, K. , Desulfurization of Stack Gas, Resources Research Institute,
Kawaguchi, Saitama, 1966.
608.* Tracer, The Decomposition Behavior of ZnSOg- 5/2 HgO. 1-9,
30 September 1968.
-------
609. United Engineers and Constructors, Inc. , Pilot Plant Study, Removal
of Sulfur Dioxide From Boiler Flue Gas, January 1964.
610. Wahnschaffe, E., Zur Entschwefelung von Rauchgasen nach dem
Dolomit-Verfahren (The Desulfurization of Flue Gases by the Dolomite
Process), VIK-Reports, 155,20-37. August 1964. ~~"~"
GOVERNMENT PUBLICATIONS
611.* Air Pollution, 1967, Air Quality Act, U.S. 90th Congress, 1st Session,
S-780, Part 4, May 1967.
612. Air Pollution Compacts, U.S. 90th Congress, 2nd Session, Parti -
Hearings before the subcommittee on air and water pollution, com-
mencing 27 February 1968, S-2350, S.J. Res. 95, S. 470. Consider-
ation of several proposed air pollution control compacts involving
West Virginia, Ohio, Illinois, Indiana, and Mid-Atlantic area.
*
613. Air Pollution Engineering Manual, Air Pollution Control District,
County of Los Angeles, U.S. Department of Health, Education, and
Welfare, 1967.
614. Air Quality Criteria for Sulfur Oxides. U.S. Dept. of Health, Education,
and Welfare, March 1967.
615. Berk, A. A. and Burdick, L. R., A Method of Test for SO, and SO, JIL
Flue Gases, Bureau of Mines, RI 4618, 1950.
616. Bienstock, D. , Brunn, L. M. , Murphy, E. M. and Benson, H. E.,
Sulfur Dioxide -- Its Chemistry and Removal From Industrial Waste
Gases, Bureau of Mines, 1C 7836, 1958. "*""""
617. Bienstock, D., Field, J. H. and Myers, J. G. , Process Development
in Removing Sulfur Dioxide From Hot Flue Gases, Bureau of Mines,
RI 5737, 1961.
618. Bienstock, D., Field, J. H. and Myers, J. G., Process Development
in Removing Sulfur Dioxide From Hot Flue Gases. 3. Pilot Plant Study
of the Alkalized Alumina System for SO- Removal, Bureau of Mines,
RI 7021, rJuly 1967.
-------
nC
619. Cinquegrane, G.C., Kurtzrock, R.C. and McCrea, D. H. , Designing
an Alkalized Alumina Pilot Plant for Sulfur Oxides Removal) Bureau
of Mines (Presented at TMS Operating Metallurgy Conference,
Extractive Metallurgy Division Symposium on the Design of Metal
Producing Processes, 11-15 December 1967).
*
620. Dean, R. S. , Present Status of Sulfur Fixation and Plan of Investigations,
Bureau of Mines, RI 3339, 3-18, May 1937
*
621. Field, J.H. , Brunn, L.W., Haynes, W. P. , and Benson, H. E. , Cost
Estimates of Liquid Scrubbing Processes for Removing Sulfur Dioxide
From Flue Gases, Bureau of Mines, RI 5469, 1959.
#
622. Fixation of Sulfur From Smelter Smoke, Progress Reports, Metal-
lurgical Division, Bureau of Mines, RI 3339, May 1937.
623.' Gartrell, F. E. , et aL , Full Scale Study of Dispersion of Stack Gases,
A Summary Report, Tennessee Valley Authority, August 1964.
624. Harrington, R. E. , Borgwardt, E.H. and Potter, A.E. , Reactivity of
Selected Limestone and Dolomites with Sulfur Dioxide, Dept. of Health,
Education and Welfare (For presentation at the American Industrial
Hygiene Conf. , May 1967, Chicago, Illinois).
625.* Katell, S. , Wellman, P., Morel, W.C., Plants, K. D. ,and Abel, W. T. ,
An Evaluation of Processes for the Removal of SO^ From Power Plant
Flue Gases, Bureau of Mines, Morgantown Coal Research Center,
Report 64-9 (Revised), 15 January 1965.
626.* Kurtzrock, R.C., Bienstock, D. and Field, J. H. , Process Develop-
ment in Removing Sulfur Dioxide From Hot Flue Gases. 2. Laboratory
Scale Pulverized Coal Fired J?'urnace, Bureau ol Mines, RI 6307, 19b3.
b27.* Kurtzrock, R.C., McCrea, D. H. and Cinquegrane, G. C. , Designing an
Alkalized Alumina Pilot Plant for Sulfur Oxides Removal, Bureau of
Mines (Presented at TMS Operating Metallurgy Conference, Extractive
Metallurgy Division Symposium on the Design of Metal Producing
Processes, 11-15 December 1967).
-------
628. Leaver, E. S. and Thurston, R. V., Ferric Sulphate and Sulphuric Acid
From Sulphur Dioxide and Air, Bureau of Mines, Ri 2556, December 1923
629. Loveless, A.H., Production of Liquid Sulphur Dioxide at I. G. Farben
Fabrik, Wolf en, British Intelligence Objectives Sub-Committee
i
630. Marks, G. W. and Ambrose, P.M., Recovery of Sulfur in Solid Com-
ounds by the Addition of Ammonia and Water Vapor to Smelter Gas,
ureau of Mines, RI 3339, 31-40, May 1937. :
631. Marks, G. W. and Ambrose, P.M., Diethylene Triamine and Other
Amines as Agents for the Recovery of Sulfur Dioxide, Bureau of Mines,
RI 3339, 41-46, May 1937.
632. • Martin, D. A. and Brantley, F. E., Selective Adsorption andJRecovery
of Sulfun Dioxide From Industrial Gases by Using Synthetic i&eolitfes,
Bureau of Mines, RI 6321, 1963.
633. Methods for Controlling Hydrogen Sulfide and Sulfur Dioxide Gases of
Refinery and Sulfur Recovery Plants, Los Angeles County APCD,
January 1966.
*
634. Roberson, A.H. and Marks, G. W., Fixation of Sulphur From Smelter
Smoke, Progress Reports - Metallurgical Division, Bureau of Mines,
RI 3415, October 1938.
635. St. Clair, H. W., Vapor Pressure and Thermodynamic Properties of
Ammonium Sulfites, Bureau of Mines, RI 3339, 19-30, May 1937.
*
636. Smith, J. R. , Hultz, J.A. and Orning, A. A., Sampling and Analysis
of Flue Gas for Oxides of Sulfur and Nitrogen, Bureau of Mines, RiTTlOS.
637. Smith, W.S. and Gruber, C.W. , Atmospheric Emissions From Coal
Combustion, An Inventory Guide. PHS Pub. No. 999-AP-Z4, i$66.
638. Spaite, P. W. , Reduction of Ambient Air Concentrations of Sulfur Oxides
Present and Future Prospects, USPHS, Presentation at National dont'er-
ence on Air Pollution, Washington, D. C. , Paper No. B-9, 12-14 Decemto
1966.
639. Sulfur Dioxide -- Its Chemistry and Removal From Industrial Waste
1958.
640. Sulfur Dioxide Removal From Power Plant Stack Gas; Conceptual Desi
and Cost Study. Sorption by Limestone or Lime; Dry Process. TVA,
-------
641. Sulfur Oxides and Other Sulfur Compounds, A bibliography with
abstracts, Dept. Health, Education & Welfare, PHS Publ. 1093,
1965.
»42. The Removal of Sulphur Gases From Smelter Fumes, Ontario
Research Foundation, Baptist Johnson, Publisher, Toronto, Canada,
1949.
643. The Third National Conference on Air Pollution. NCAPC, Washington,
D. C. , 12-14 December 1966.
*
644. Wartman, F.S. , Oxidation of Ammonium Sulfite Solution, Bureau of
Mines, RI 3339, 47-51, May 1937.
645. Wells, A. E. , Thiogen Process, Bureau of Mines, Bulletin 133, 1917.
BOOKS
646. Air Pollution Abatement Manual. Washington, D. C., Manufacturing
Chemists Assoc. , Inc. , 1952.
647. Aries, R.S. and Newton, R. D. , Chemical Engineering Cost Estimation.
New York, McGraw-Hill, 1955.
648. Audneth, L. F. and Ogg, B.A., The Chemistry of Hydrazine, New York,
Wiley and Sons, 1951.
649. Bauman, H. C. , Fundamentals of Cost Engineering in the Chemical
Industry, New York, Reinhold, 1964.
650. Boynton, R. S. , Chemistry and Technology of Lime and Limestone,
New York, Interscience Publishers, 1966.
-------
651. Chemical Economics Handbook, Menlo Park, Cal., Stanford Research
Institute, " Ammonia -Salient Statistics, " 703.4330A and B, 1966.
652. Clark, C.C., Hydrazine, Baltimore, Md., Mathieson Chemical Corp.,
1953.
653. Faith, W. L., Keyes, O. B. and Clark, R. L., Industrial Chemicals,
New York, Wiley and Sons, 1965. ~~
654. Kirk, R. E. and Othmer, D.F. (eds.), Encyclopedia of Chemical
Technology, New York, Interscience, 1951. "Ga s Cleaning and
Purification -- Removal of Sulfur Dioxide, " 7, 103-4.
655. Ibid., Second Edition, 1963. "Ammonium Sulfate, "£, 329-30.
656. Ibid., Second Edition, 1964. "Calcium Sulfate, " 4, 14-27.
657. Ibid., Second Edition, 1966. "Nitrogen Fertilizers, " 9, 51-78.
658. Ibid., Second Edition, 1966. "Gas Cleaning, " l±, 329-52.
659. Ibid., Seeond Edition, 1966. "Hydrazine and its Derivatives, "
11, 164-96.
670. Kohl, A. L. and Riesenfeld, F.C., Gas Purification, New York,
McGraw-Hill, I960.
671. Latimer, W.M., Oxidation Potentials, New York, Prentice-Hall, 1952.
672. Magill, P. L. , Holden, F. R. and Ackley, C. (eds.), Air Pollution
Handbook, New York, McGraw-Hill, 1956.
673. Mallette, F.S. (ed.), Problems and Control of Air Pollution, New York,
Reinhold, 1955. [ ~~~
674. Meetham, A.R. , Atmospheric Pollution - Its Origins and Prevention,
Pergamon Press, London, 1952.
675. Nonhebel, G. (ed.), Gas Purification Processes, London, Geroge
-------
676. Perry, R.H. , Chilton, C.H. and Kirkpatrick, S. D., (eds.), Chemical
Engineers' Handbook, New York, McGraw-Hill, 4th ed,, "Gas Absorption, "
14, 1-40,1963.
677. Robinson, C.£i. , The Recovery of Vapors. New York, Reinhold,
678. Schroeter, L. C. , Sulfur Dioxide, London, Pergamon Press, 1966.
679. Scorer, R.S. . Air Pollution, London, Pergamon Press, 1968.
680. Spengler, G. and Michalczyk, G. , Die Schwef el -oxide in Rauchgaaen
und der Atmosphare (mit umfassendem Literaturverzeichnis) [Sulfur
Oxides in Flue Gases and in the Atmosphere (with Comprehensive
Bibliography)]. Dusseldorf, VDI-Verlag GmbH. , T964.
681. Stern, A. C. (ed.), Air Pollution, New York, 3 Vols, Academic Press,
Sec. ed, 1968.
682. Strauss, W. , Industrial Gas Cleaning, London, Pergamon Press, 1966.
683. Sulfur Dioxide Technical Handbook, Atlanta, Ga. , Tennessee Corp.
THESES AND DISSERTATIONS
684. Coke, J. R. , "The Removal of Sulphur Oxides From Waste Gases by a
Dry Method, " Doctoral Thesis, University of Sheffield, England, Dept.
of Fuel Technology, May I960.
685. Coughanowr, D. R. , "Oxidation of Sulfur Dioxide in Drops, " Ph. D. Thesis
in Chemical Engineering, University of Illinois, 1956 (University
Microfilms, Inc., Ann Arbor, Mich.).
686. Galeano, S. F. , "Sulfur Dioxide Removal and Recovery in the Pulp Mill
Industry, " Doctoral Dissertation, University of Florida, August 1966.
687. Krause, F. E. , "The Reaction of Sulfur Dioxide and Oxygen m Aqueous
Solution Containing Manganous Sulfate as a Catalyst, " M.S. Ihcsis in
Chemical Engineering, Purdue University, 1959.
-------
688. Terraglio, F. P., "Laboratory Evaluation of Methods for Sulfur
Dioxide," M.S. Thesis, Rutgers University, 1962.
689. West Wm. E., Jr., "Evaluation of Sulfur Dioxide Recovery Processes*"
M.S. Thesis, University of Illinois, 1953.
-------
AUTHOR INDEX
(excluding principal authors)*
Abel, W. T., 625
Ackley, C., 672
Akbar, M. , 179
Ambrose, P. M. , 630-1
Amsler, R. L. , 28
Andres, A. S., 440-1
Aniol, S., 324
Annenkov, V. A. , 242
Aristov, G. E. , 93
Atsukawa, M. , 438
Avdeeva, A. V. , 466-8
Babaev, E. V. , 16
Bacia, W., 42
Balarev, K. , 301
Bang, Do Hoon, 244
Barber, F. R., 168, 293
Barney, J. E. , 24
Bartunek, A. , 572
Bauer, E. R. , Jr. , 28
Bellen, A. , 286
Benson, H. E. , 30, 141-2, 616
Berard, P. A., 364
Berk, J. M. , 186
Bettelheim, J. , 252-3, 532-3, 536
Bienstock, D., 626
Bistron, S. , 324
Blankmeyer, H. C. , 224
Bohac, J., 572
Borgwardt, E. H., 624
Boving, 539
Box, P. L. , 238
Boyadjiev, H. , 124-6
Brantley, F. E., 632
Bregeault, J. M. , 318
Bretsznajder, S., 505
Brice, D. B., 103-4
Brooks, A. F., 120
Brunn, L. M., 141-2, 616
Burdick, L. R. , 615
Bushmelev, V. A. , 287
Cada, V- , 389
Carpenter, S. B. , 161-2
Carson, J. E., 304
Cauldfield, J. G. L., 205
Chapin, J. H., 595
Chass, R. L. , 164
Chernov, E. N. , 469-70
Chertkov, B. A., 9, 332
Chilton, C. H. , 13, 676
Chlumsky, F. , 25, 536
Chueh, Ping Lin, 4
Cinquegrane, G. C. , 627
Cirey, 514
Clarke, A. J. , 412
Principal authors are listed alphabetically in the bibliography.
-------
AUTHOR INDEX
Cole, R. J., 234
Cottrell, 539
Coughanow*. D. R., 215
Crainpton, E. W., 243
Creanga, L., 346
Cylmer, A. B., 432
Davis, J. L., 455
Decry, R. F., 119
Delmarcel, G., 144
Derka, J.. 508
CKckerson, R. C., 407
Dickey, S. W., 336
Djega-Mariadassow, G., 318
Doyle, G. J., 361
Driskell, J. C., 424
Dvorak. K., 37
Eaven, A., 309
CHrnbv, V. T., 172
Efimova, T. F., 442
Elenkow, D., 301
Fayer, S. M., 247
Fernandez, C. I., 440-1,
Fernelius, W. C., 45
Field, J. H., 29-32, 334, 479,
617-8, 626
Fischer, F., 527
Fitt, T. C., 145
Fletcher, A. W., 114
Frantitafc, F., 508
Frieling, G., 343-4
Fukui, S., 193, 438
Gaeke, G. C., 452
Gall, G. F., 601
Gandon, L., 543
Cans, S. N.. 264
Gantz, R., 43
GavlovBkaya, S. S., 340
Gazaryan, S. A., 289-90
George, R. E., 54
Gerstlc, R. W., 583
Gieasen, J. A. van der., 179
Giguere, P. A., 137
Gleason, T. G., 257
Gofman, M. S., 442
Goldman, H., 506
Gorbachev, B. I.. 265
Gray, F. J., 424
Grigoryan, G. O., 289-90
Grigoryan, M. M., 290
Grossinsky, O., 502
Grossley, H. E., 335
Gruber, C. W., 637
Haar, L. W. ter., 179
Hahn, E., 434
Hala, E., 37
-------
AUTHOR INDEX
Hamming, W. J., 286
Han, S. T. , 399, 455
Harding, C. I., 156
Harkins, J. , 375-6
Harrington, R. E. , 348
Hasegawa, H., 542
Haynes, W- P., 141-2
Helfrich, E. , 435
Hene, W. , 360
Heredy, L. A. , 174
Higashi, M. , 438
Himmelblau, D. M. , 449
Hiser, L. L. , 377
Holden, F. R. , 672
Hultz, J. A. , 636
Ikonopisov, S. , 127-8
Isaeva, N. M. , 52, 287
Jackson, A. , 462
Jaeger, L., 568
Jara, V. , 535
Johnstone, H. F. , 165
Jonakin, J., 300, 341, 602
Jones, T. M. , 424
Juntgen, H. , 121, 262
Kamack, H. J. , 269
Kamet, K. , 193
Kapustova, J., 561
Karasev, A. G., 310
Karbanov, S. V., 351
Karzhavin, V. A., 466-8
Kas'yan, D. T., 231
Katell, S. , 31
Kawecki, W. , 42-4, 505
Kellogg, H. H. , 200
Kenroe, W. H. , 330
Ketov, A. N., 331
Keyes, D. B. , 216
Kikuchi, S. , 418
Kirkpatrick, S. D. , 676
Kirkpatrick, W. S., 274-5
Kironuma, H. , 249
Kitts, A. P., 340
Kleinschmidt, R. V-, 217
Klempt, W., 487, 502
Klimecek, R., 25, 508, 521, 533,
568, 571
Klohr, J. W. , 391-2
Kopylev, B. A. , 351
Kordik, E. , 533, 571
Kostadinov, N. , 201
Kotowska, W. , 43-4
Kramer, G. D. , 407
Kraus, H. , 527
Krause, F. E., 107
Krechemov, T. T., 467
Krustev, I. , 126
Kubel, K. , 146
Kucheryavyi, V. I., 170
Kurtzrock, R. C., 143, 619
Kuznetsov, I. E. , 158-60
-------
AUTHOR INDEX
Larson, G. P. , 579
Lengyel-Szabo, G. , 263
Lenher, S., 425
Leont'eva, L. S., 231
Lepper, G. H., 149
Leppla, P. W-, 218
Lewis, H. C., 371
Leykin, L. I., 160
Little John, R. F., 106
Litvinenko, I. I. , 172
Logan, L., 198
Ludwig, J. H., 367-8
Lukacs, J., 372
Lundberg, L. A. , 327
\
Lyons, C. J., 327
Maksimov, V. F, , 52
Mandersloot, W. G. , 352
Manganelli, R. M., 428
Manning, R. N., 267
Markant, H. P., 499
Marks, G. W., 634
Martin, A., 168
Martin, J. R. , 602
Matsumoto, K., 15, 475
Mattia, M. M. , 365
Matty, R. E. , 598
McCarthy, J. L., 327, 382
McConnell, F. J. , 460
McCrea, D. H. , 143, 619, 627
McLLroy, R. A. , 598
McKee, H. C. , 377
McKenzie, D. E., 174
Meisiel, H., 528
Michaels, A., 579
Michalczyk, G. , 680
Miller, L. A. , 429, 569
Miller, P., 448
Miller, R., 488
Mirumyan, R. L., 290
Mitchell, R. F. » 496
Mitra, A. K., 323
Molchanova, N. I., 16
Moll, A. J., 219
Moore, D. J., 280
Morel, W. C. , 625
Mulvihill, J. W., 621
Murphy, E. M., 616
Myers, J. G., 32, 617-8, 621
Nagai, S., 12
Namba, Y., 422
Nankov, N., 127-8
Neville - Jones, D. , 575
Newall, H. E., 303
Newton, R. D., 646
Nicolet, S. , 46
Nicol'skaya, Y. P., 468
Nieuwenhuizen, J. K. , 179
Nikolaev, A. M., 378
Nishimoto, Y., 15, 475
Nonhebel, G. , 328-9
-------
AUTHOR INDEX
Novella, E. C. , 388
Novikov, A. I. , 339
I, B. A. , 647
Olden, J. M. E. , 49
Ormng, A. A. , 583, 636
Othmer, D. F. , 654
Paillard, H. , 46
Papyan, G. S. , 290
Patrick, W. A. , 295
Payne, J. W. , 105
Pearce, S. L. , 192
Pearson, J. L., 500, 550
Pekareva, T. I. , 94
Peters, W. , 229
Peterson, D. G. , 138
Phillips, N. D. , 292
Phillips, W. M. 579
Pigford, R. L. , 453, 595
Pike D. E. , 459
Pmaev, V. A. , 553
Pitelma, N. P. , 339
Plants, K. D. , 31, 233, 625
Plummer, A. W. , 118
Podgaiko, V- V. , 159
Polem, J. , 561
Poll, A. , 110-1
Polhtt, A. , 192
Ponyankov, B., 436
Popov, D. M. , 265
Porr, A. , 516
Porter, M. A. , 27
Potter, A. E. , 624
Pukhna, D. L. , 93-5
Read, H. J. , 224
Redferan, M. W. , 116
Rees, R. L. , 191-2
Remauer, T. V. , 377
Renwick, C. W. , 492
Reverben, A. , 139
Riesenfeld, F. C. , 255, 670
Rispler, L. , 370
Roesner, G. , 450
Rossano, A. T. , 285
Ryabenko, I. M. , 563
Ryan, T. G. , 228
Safin, R. Sh. , 310
Saline, L. E. , 177
Schauer, V., 37
Schwartz, C. H. , 583
Scribner, A. K. , 194
Sesenbaugh, J. D. , 138, 279
Shah, I. S. , 292
Shannon, P. T. , 4
Shlifer, V. A. , 160
Shligerskii, A. S. , 239
Shroeder, M. , 183
Shulman, H. L. , 100
-------
AUTHOR INDEX
Shutkin, G. A. , 242
Shutko, F. W., 341
Sieth, J., 187
Simons en, R. N., 261
Singer, J. C., 602
Singh, A. D., 220-1, 594
Skrivanek, J. , 253, 532
Slodyk, T., 2
Smith, N., 374
Snell, H. A., 393
Snowball, A. F. , 184
Solbett, J. M. , 202
Sosekina, G. V., 339
Spaite, P. W., 182, 282, 348
Spurr, G. , 280
Stary, M. , 402
Steigcrwald, B. J. , 283, 368
Stone, M. H. , 306
Streltsov, V. V., 6
Striplin, M. M., Jr. , 448
Strnad, M. , 25, 536, 571
Strom, G. H., 304
Swearingen, L. E. , 362
Synoradski, Z. , 325
Szaraware, J., 325
Takenouchi, H. , 12
Tarantino, S. , 99
Tarat, E. Ya. , 351
Taube, H., 135
Taylor, A. , 530
Teodorescu, C., 346
Thodos, G,, 117
Thomas, F. W., 161-2
Thorogood, A. L. , 575
Thurston, R. V., 628
Tomany, J. P., 343-4
Tseittin, A. N. , 172
Tuerkoelmez, S., 254
Ulander, P. H. N., 328-9
Umbach, H., 502
Uwanishi, G., 249
Valea, I., 19
Vivian, J. E., 454
Wainwright, H, W. , 621
Walker, W. J. S., 110-1
Walter, D. F. , 188
Wanjukow, W. , 195
Watabe, H., 542
Wellman, P. , 625
West, F. B. , 327
West, W. E. , 222, 490
Whiddon, O. D., 341
Whitney, R. P. , 454
Winsche, W. F. , 223
Wolf, F. , 153
Wood, C. W., 238
-------
AUTHOR INDEX
Yamada, H. , 438 Zalogin, W. G. , 562
Yamada, N. , 302 Zapryanova, A. , 436
Yosiu, S. J. , 174 Zawadski, E. A., 169
Yu, Sun-Nien, 4 Zhavoronkov, N. M. , 378
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APPENDIX A
THEORETICAL CHEMICAL EQUILIBRIA CONSIDERATIONS
The following three equilibria govern the solubility of SO, in aqueous
media:
the equilibrium between the gas and unionized dissolved SCL,
H,SO. = SO, (gas) + H,0
2 3 Z *
so
K,
(H2S03)
(A-l)
the first lonization of aqueous
HS0
K = (H } (HS°3
(H2S03)
the second lonization of aqueous SO,.
HS03~ = H" + S03
Let s = the total concentration of SO3 in solution, i. e. ,
(SO =)
(HS03~)
s = (H2S03) + (HS03~) + (S03=)
(A -2)
(A-3)
(A -4)
Equations A-l to A-3 can then be combined to yield Equation A-5 which relates
the partial pressure of SO, over the solution, the total amount of SO2 in
£
solution, and the hydrogen ion concentration.
S0
(H+)Z
CA-5,
An examination of Equation A-5 indicates that the partial pressure of
in equilibrium with the solution can be reduced to any amount desired
-------
by reducing either the hydrogen ion concentration or the amount of SOg in
the solution. Johns tone22 gives the following values for Kj, K£l and K}
at 25°C:
, Kj = 0. 795 atm/mole/1000 g HgO
KI = 0.013 mole/1000 g H2O
K3 = 10"7 mole/1000 g H2O.
If the assumption is made that any scrubbing solution will have
1 mole /liter of SO, dissolved, then the partial pressure of SO2 will be
less than 150 x 10 atm in any solution in which the pH is higher than 5. 6.
These constants are temperature dependent and somewhat composition
dependent.
The reader will note that many of the considerations which follow
are also attributable to Johnstone. A charge-balance equation is necessary
to relate the hydrogen ion concentration to that of other species present.
Let M. be the concentration of a charged ion species, i, and let Z^ be the
charge (either a positive or a negative integer). All species other than
H , HSO,~, OH~, and SO ~ are considered. Then the net charge in the
solution must be zero and Equation A-6 follows:
(H*) + ? Zi Mj = (HSO3~) + 2(SO3=) + (OH").
The use of Equations A-2 and A-3 and the definition of s with Equation A.-6
result in Equation A-7.
Substitution of Equations A-2, A-3, and A-6 in Equation A-4 and
rearranging gives:
' | + £ Z.M.-(OHT]
s =
Combining Equations A -5 and A-7 results in:
Kj (H+)2 £ (H^) + £ Z.M. - (OH'fJ
-------
The hydrogen ion concentration can be eliminated from Equations A-7 and A-8
to provide an expression for a, the sulfur dioxide solubility, and P0_ , the
rf" SO-
partial pressure of SO2> in terms of ^ Z^ M^, the other constituents present.
As an example, consider the case in which ammonia is the only other
species present in total concentration, C, and in acid solution of pH 4-7. Then
(NH4+) = £ ZL Mx = C and
Equation A-7 thus becomes: (H+) = ?A?ff K
-
K, K- ._ c -..2
and Equation A-8 becomes: J "
K1K3
-w — can be combined in one constant, M, which is a function of temperature and
^ 224
depends slightly on concentration. Johnstone gives a measured value for M in
ammonia of 20 x 10 atm/mole /liter H-O at 25°C which agrees fairly well with
L <*
the calculated K,K.,/K0 of 6. 1 x 10 atm/mole /liter H,O assuming ideal solution.
1 j £. £
Johnstone has also developed a series of equations to represent the results
Q YD
of the measurement of SO- pressure over concentrated alkaline solutions.
P -
Pso2 '
where: Pc/^ = SO, pressure in mm
JjvJ^ £•
S = Total SO- in solution, moles/ 100 moles H2O
C = Total moles of base in solution, moles/ 100 moles H-.O
F (T) = 0. 0343 for ammonia solutions at 50 C
F (T) = 0. 0233 for sodium hydroxide solutions at 50 C.
-------
These numbers were used in deriving the values for the ammonia and sodiunf
hydroxide scrubber systems shown in Table 2 of the text.
The question arises as to whether carbon dioxide will displace SO, from
an aqueous solution and decrease the capacity of the solution for absorbing SO2<
The equilibrium involved in an aqueous solution in which all the constituents
are soluble is:
CO, + HSO " = SO, + HCO "
2323
671
Using the thermodynamic functions from Latimer (see Table A-l), the free
energy change of this reaction is found to be +8.2 kcal at 25 C and the entropy
change is +5. 04 entropy units. Since the entropy change is positive, the reaction
is driven to the right by raising the temperature. As a first approximation for the
temperature correction in raising the system to 50°C, the heat capacities of the
substances represented on both sides of the equation can be assumed to be equal.
Since dAF°/dT = - AS0, AF° at 50° = AF° at 25° - 25 AS° = 8.0 =
-RT InK.
The equilibrium constant for the reaction at 50°C is estimated to be
4 x 10"6. Then:
(HC03~)
(HS03")
At the bottom of the scrubber where P_~ =112 mm and Pe_ = 2. 28 mm, the
A 2 ***"'2
bicarbonate to bisulfate ratio is 2 x 10 .' This means that virtually none of the
aqueous base at the bottom of the scrubber is in the form of bicarbonate and is
therefore used in absorbing SO,.
The situation for slurries of basic oxides is somewhat different. For both
calcium and magnesium oxides, the 14. 7% partial pressure of carbon dioxide in
flue gases is sufficient to convert the oxide to the carbonate. Hence, the net
reaction to be considered is MCO, + SO, = MSO- + CO,, where M represents
either calcium or magnesium.
-------
TABLE A-l
THERMODYNAMIC
Formula
Ca++
CaO
Ca(OH)2
CaSO3
CaCO3
C02
co2
H2C03
HC03"
co3=
H+
OH"
H20
H20
Mg++
MgO
Mg(OH)2
MgS03
MgC03
so2
so2
HS03"
H2S03
so3=
State
aq
C
C
C
C
g
aq
aq
aq
aq
aq
aq
aq
g
aq
C
C
C
C
g
aq
aq
aq
aq
DATA ON SOME COMPOUNDS OF INTEREST
H°
kcal/mole
-129. 77
-151. 9
-235.8
-
-288.45
-94.0518
-98.69
-167 0
-165. 18
-161.63
0
-54.957
-68.317
-57.798
-110.41
-143.84
-221.00
-241.0
-266.
-70 76
-80. 86
-151.9
-145. 5
-151.9
F°
kcal/mole
-132. 18
-144.4
-214. 33
-
-269. 78
-94.2598
-92.31
-149.00
-140.31
-126.22
0
-37.595
-56.690
-54.635
-108.99
-136. 13
-199.27
-221. 2
-246
-71.79
-126.0
-128. 59
-116. 1
S°
cal/deg K
-13.2
9.5
18.2
24.2
22.2
51.061
29.0
45. 7
22. 7
-12.7
0
-2.52
16. 716
45. 106
-28.2
6.4
15.09
22. 5
15.7
59.4
26.
56.
-7.
-------
The free energy of formation pf calcium sulfite is estimated to be
-255.25 kcal/mole. This value is derived from Latimer's free energies of the
separate ions and from the solubility product of calcium sulfite. The following
table can then be derived:
F° 2
Reaction (kcal/mole) K 14. 7% (mm) CO?
a. CaCO3 + SO2 = CaSO3 + CO2 -7. 9 6. 6 x 105 1. 7 x 10~4
b. MgCO3 + S02 = MgSO3 + CO2 +2. 33 0.02 5600
c. MgCO +H20 +
+ 2 HS03"
d. CaCO +H20 + 2S02 = _Q ^ 1.4 xlO6 9x10°
Ca"*"*" + 2HSO3" + CO2
These values indicate that if the carbon dioxide pressure is higher than
_A
1. 7 x 10 mm, Reaction (d^with calcium carbonate, will occur. With magnesium
carbonate, Reaction (b) will not occur unless the SO- pressure is higher than
tf i
5600 mm, but Reaction (c) will occur if the carbon dioxide pressure is higher than
5 x 10" mm.
At the top of the scrubber, where the SO, pressure is 0. 1 mm (0. 015%),
these numbers indicate that SO, will be absorbed by calcium carbonate with the
formation, in solution, of calcium bisulfite. A small change in the values of the
equilibrium constants would suggest that calcium sulfite would be precipitated.
Magnesium carbonate would be expected to absorb SO- with the liberation
of carbon dioxide and there is small likelihood of the precipitation of magnesium
sulfite. These results are consistent with the fact that calcium sulfite is much
less soluble than magnesium sulfite. A temperature correction has not been
applied. The uncertainty in the data is much larger than the factor of about
0. 2 kcal/mole correction to the free energy which is needed for the temperature
correction.
-------
Calcium or magnesium oxide slurries in amounts in
excess of the amount of SO- present will absorb
carbon dioxide and be converted to the carbonates
Both calcium and magnesium carbonate are basic
enough to absorb SC>2 at the 0.015% level with
calcium having the larger driving potential.
There is a moderate possibility of precipitating
calcium sulfite, even in fairly dilute solutions.
There is nothing in the available thermodynamic
data to indicate that calcium or magnesium oxide
slurries are incapable of reducing SO- partial
pressures to 0.015%.
-------
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-------
APPENDIX B
CONVERSION OF GASEOUS SULFUR DIOXIDE TO
MARKETABLE PRODUCTS: COST ESTIMATES
GENERAL
Many of the processes discussed in this report yield sulfur dioxide as a
by-product. For each of these systems, both the capital and operating costs
required to liquefy the sulfur dioxide were included in the total costs.
The potential quantities of sulfur dioxide produced by recovery from flue
gas are so large that it is unlikely that the market for liquid SO, could absorb
£t
this production. Since approximately 80% of the natural sulfur mined annually
is converted to sulfuric acid, it is logical to assume that the optimum usage
for sulfur dioxide recovered from flue gas should be in the production of
sulfuric acid. The following three approaches are available for accomplishing
this goal:
• Liquefy the sulfur dioxide for shipment to sulfuric
acid producers.
• Produce sulfuric acid from the recovered gaseous
sulfur dioxide at the power plant site.
• Reduce the gaseous sulfur dioxide to elemental
sulfur, which in turn can be converted to sulfuric
acid as warranted.
The selection of the specific route would depend on the sulfuric acid market,
on the existing sulfuric acid producers in the locality of each power plant, and on
relative shipping costs.
Order-of-magnitude capital and operating cost estimates have been prepared
for each of the three approaches. These are presented below.
LIQUEFACTION OF SULFUR DIOXIDE
The theoretical quantity of sulfur dioxide produced in most of the processes
which yield SO as a by-product is 34, 610 tons per year. This is based on the
standards of the 120-megawatt power plant used in Phase I of this study.
Capital Costs
As mentioned in the general discussion above, these costs have been
included in the total capital cost of each process, where applicable. The purchased
-------
equipment cost, taken from the Bureau of Mines study, is $102,000. This pro-
vides a fixed capital cost of $483, 500 and a total investment including working
capital of $531,400. The capital requirements are: $4.43 per kw capacity,
and $15. 25 per ton SO. produced. All of these values are deductible from the
specific process capital costs to obtain the net process costs.
Operating Costs
The operating costs have been summarized in Table B-l. Purchased
raw materials are not required. Direct labor has been estimated at one man per
shift. The supervision requirement has not been increased, since it would be
provided from the SO, recovery system operation. Utilities requirements are
138, 000 kwh electricity and 623, 000 M gal circulating cooling water.
SULFURIC ACID PRODUCTION
The sulfur dioxide recovered would yield approximately 50,000 tons of 1009
sulfuric acid per year. The most economical production cost would be in a
system where the gaseous SO. could be converted directly to sulfuric acid at the
to
power plant site.
Capital Costs
Capital investment required for a contact process sulfuric acid plant
varies from $12, 500 per ton for a 100-ton per day plant to $10, 000 per ton for a
240-ton per day plant. A 150-ton per day acid plant at $12, 500 per ton would
cost $1, 875, 000. Adding working capital, the capital requirement is $2, 062,500,
or $17.20 per kw capacity. This corresponds to $59. 60 per ton SO. converted
and $41. 25 per ton H.SO, produced.
Operating Costs
653
The reference cited lists the following requirements per ton of
100% sulfuric acid produced:
Labor - 0. 64 man hour
Electricity - 5 kwh
Water - 4,000 gal
Steam - 200 Ib
Air - 250,000 cu. ft.
The direct labor requirement is 32,000 man hours, or 16 men (4 per
shift). One foreman per shift has been charged to the system.
-------
TABLE B-l
LIQUEFACTION OF SO2*: OPERATING COST ESTIMATE SUMMARY
Fixed Capital Cost: $483, 500
ITEM TOTAL $
1. Raw Materials fc Chemicals
*
34,610 tons per year
2. Direct Labor 24. OOP 13. QQ
3. Supervision _ __-
4. Maintenance, 5% of Fixed Capital 24,200 13. 96
5. Supplies, 15% of Maintenance 3, 600 2. 08
6. Utilities 31.800 18.41
7. Other - «_^«_
8. TOTAL DIRECT COST 83,600 48.35
9. Payroll Burden, 20% of 2 & 3 500 0.29
10. Plant Overhead, 50% of 2, 3, 4 & 5 25,900 15.00
11. Pack & Ship - -
1Z. Waste Disposal - -
13. Other - _.
14. TOTAL INDIRECT COST 26.400 15.29
15. Depreciation, 10 % Fixed Capital/Yr 48,400 27.96
16. Taxes, 2% of Fixed Capital 9.700 - 5^62-
17. Insurance, 1% of Fixed Capital 4.800 2-78
18. Other _ = - - = -
19. TOTAL FIXED COST 62,900 36.36
20. TOTAL OPERATING COST 172,900 IOQ.QQ
21. COST: $/Ton of Coal 0.36
22. Mill/kwh 0. 18
$/Ton SO? 5.00
$/Long Ton S Equivalent: 11.20
-------
1 The utility requirements are:
Power - 250,000 kwh per year
Water - 200, 000 M gal per year
Steam - 10, 000 M Ib per year
Table B-2 summarizes the operating costs. The operating cost
of $12. 97 per ton of 100% sulfuric acid may have some significance since the cost
of purchased sulfur to produce one ton of 100% sulfuric acid in a conventional
plant is $11. 40 (0. 3 long ton S at $38 /long ton). Addition of this raw material
cost would increase the operating cost to $24. 37 in a conventional sulfuric acid
plant. The manufacturing cost in large sulfuric acid plants is undoubtedly less
than $24. 37 per ton 100% H-SO,. This cost comparison, however, illustrates
the possible cost reductions achievable by eliminating the raw material cost
in sulfuric acid production via the contact process.
73
A recent announcement indicates that sulfuric acid could be
produced from gypsum in a large plant for $16 per ton. Investment for a
1500-ton-per-day sulfuric acid plant (and a roughly equivalent amount of cement
by-product) would be about $30 million.
REDUCTION OF SULFUR DIOXIDE TO SULFUR
Reference is made in the literature to the Guggenheim process, a
method in which the SO, is reduced to sulfur. The SO- is mixed with 25 percent
air, preheated to 400°C, and passed through a coke bed which is maintained
incandescent (800°C) by the reaction between coke and air. The gases pass
through the reduction chamber at such a rate that a substantial amount of
remains. The exit gases, which contain sulfur, carbonyl sulfide (COS),
carbon disulfide (CS-), and SO, are passed to a second chamber containing
a pumice catalyst; the carbon -containing compounds react with SO, to form sulfu
and CO. the sulfur is then condensed and the exit gases are returned to the
absorption system.
Capital Costs
No recent cost data were found in the literature. The reference
-------
TABLE B-2
PRODUCTION OF SULFURIC ACID FROM GASEOUS SO *:
OPERATING COST ESTIMATE SUMMARY
Fixed Capital Cost: $1, 875, 000
ITEM TOTAL $
1. Raw Materials & Chemicals
50,000 short tons SO- per year
B-5
2. Direct Labor 96,000 14.81
3. Supervision 31,200 4.81
4. Maintenance, 5% of Fixed Capital 93,800 14.47
5. Supplies, 15% of Maintenance 14, 100 2. 17
6. Utilities 26,500 4.08
7. Other
8. TOTAL DIRECT COST 261.600 40.34
9. Payroll Burden, 20% of 2 & 3 25,400 3.92
10. Plant Overhead, 50% of 2, 3, 4 fe 5 117. 60fr 18. 14
11. Pack 8c Ship - -
12. Waste Disposal - -
13. Other
14. TOTAL INDIRECT COST 143.000 22.06
15. Depreciation, 10 % Fixed Capital/Yr 187.500 28.92
16. Taxes, 2% of Fixed Capita) 37.500 5.78
17. Insurance, 1% of Fixed Capital 18.800 2. 90
18. Other - -
19. TOTAL FIXED COST 243.800 37.60
20. TOTAL OPERATING COST 648.400 100.00
21. COST: $/Ton of Coal 1. 36
22. Mill/kwh 0.68
-------
The Marshall and Stevens Index was used to convert this 1937 cost to present day
costs; the C. £: Index was not used since it was not developed until 1^47. Since
the slope of the M & S Index for the post 1947 period is practically the same as
the C. E. Index, its use is considered to lead to costs which are comparable to
other factored costs in this study. The M & S Index indicates a three-fold
increase in capital costs between 1937 and 1967. On this basis, it was assumed
that the capital cost for the SO, reduction system would be $45 per ton of sulfur
to
per year, or approximately $780, 000 in fixed capital cost. Adding 10% for
working capital, the unit costs are $7. 15 per kw capacity, a value equivalent to
$24. 80 per ton SO, converted or $55. 50 per long ton S produced.
Operating Costs
Table B-3 summarizes the operating costs. The raw material cost ,
is for coke at the rate of 0. 75 ton of coke per ton sulfur. On this basis, approxi-
m ately 13, 000 tons of coke per year are required.
A direct labor requirement of one man per shift is considered
adequate. Supervision can be provided from the SO, absorption plant operation.
The utility requirements are as follows:
Power (for preheating SO, gas): 2 x 10 BTU per year
b
Circulating Water (for cooling
the sulfur and carbon
monoxide and condensation
and cooling of sulfur to 300°C): 36, 300 M gal per year
-------
TABLE B-3
REDUCTION OF GASEOUS SO2 TO ELEMENTAL SULFUR*:
OPERATING COST ESTIMATE SUMMARY
Fixed Capital Cost: $778, 500
ITEM TOTAL $ %
1. Raw Materials & Chemicals 247,000 51.26
2. Direct Labor 24.000 4.98
3. Supervision - -
4. Maintenance, 5% of Fixed Capital 38,900 8.07
5. Supplies, 15% of Maintenance 5,800 1.20
6. Utilities 30.000 6.23
7. Other - -
••••••^•••••B aVMHMBHBn
8. TOTAL DIRECT COST 345, 700 71.74
9. Payroll Burden, 20% of 2 fc 3 500 0. 10
10. Plant Overhead, 50% of 2, 3, 4 k 5 34.400 7.14
11. Pack b Ship - -
12. Waste Disposal - -
13. Other - -
14. TOTAL INDIRECT COST 34.900 7.24
15. Depreciation, 10 % Fixed Capital/Yr 77.900 16. 16
16. Taxes, 2% of Fixed Capital 15.600 3.24
17. Insurance, 1% of Fixed Capital 7.800 1.62
18. Other - -
19. TOTAL FIXED COST 101.300 21.02
20. TOTAL OPERATING COST 481.900 100.00
21. COST: $/Ton of Coal 1.Q2
22. Mill/kwh 0.51
$/L ton S 31. 19
15,450 L. tons sulfur per year
-------
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-------
APPENDIX C-l
Acknowledgments
The contributions of the following equipment companies are
acknowledged.
Allis Chalmers
Bird Machine Company
Buell Engineering Company, Inc.
Buffalo Forge Company
The Ceilcote Company
Chemineer, Inc.
Chicago Bridge and Iron Company
Copolymer Corporation
De Laval Separator Company
Dorr-Oliver, Inc.
The Eimco Corporation
FMC Corporation - Hydrodynamics Division
Goulds Pumps, Inc.
Ingersoll-Rand Company
Joy Manufacturing Company
Parker Brothers, Inc.
Sprout Waldron and Company, Inc.
Struthers-Wells Corporation
U. S. Stoneware, Inc.
UOP Air Correction Division
-------
ZINC OXIDE PROCESS
Case 1 & 2
Equipment List
Item Cl - Flue Gas Blower (2 required)
Capacity: 1, 675, 000 ACFM @ 500 RPM
Static Pressure: 5 in. S.W.G.
Motor HP: 2250
Brake HP: 2140
Fan Diameter: 15 ft
Diffuser: 40 ft long x 21 ft exit diameter
Weight: TO.OOOlb
Price: $128,000
Item C2 - Dryer Gas Blower (1 required)
Capacity: 311, 650 ACFM @ 558 RPM
Static Pressure: 5 in. S.W.G.
Motor HP: 350
Brake HP: 330
Wheel Diameter: 7 ft 5 in.
Price: $39,200
Item C3 - SO2 Recycle Compressor (1 required)
Capacity: 1890 ACFM® 3160 RPM
Intake Pressure: 14. 7 psia
Discharge Pressure: 22.7 psia
Intake Temperature: 120°F
Discharge Temperature: 203°F
Motor HP: 100
Brake HP: 83
Price: $6863
Item El - SO2-H2O Condenser (1 required)
Type: Direct contact packed column
Size: 12 ft dia x 20 ft high steel column epoxy resin coated
Packing: 904 cu ft 3 in. stoneware, Intalox saddles
AP: 2. 4 in. S.W-G. across packing
Price: $17,000
-------
Item E2 - Heat Exchanger (1 required)
Duty: 81,000,000 Btu/hr
Area: 7650 sq ft
Size: 34 in. I. D. steel shell
Tubes: 3/4 in. O. D. x 20 ft long, type 31o SS
Tube Plates: C.S. clad with type 316 SS
Max. Oper. Pressure: 100 psig, both tube side and shell side
Price: $46, 170
Item Ml - Ash Filter (2 required)
Type: Vacuum Drum Filter, 12 ft dia. x 14 ft long,
532 sq ft rubber-covered, complete with
filtrate receiver, moisture trap, filtrate
pump and vacuum pump.
Capacity: 163 tons/day dry cake
Total Motor HP: 81. 5
Price: $55,220
Item M2 - Zinc Sulfite Centrifuge (5 required)
Type: 36 in. x 96 in. screen bowl centrifuge
Motor HP: 250
Capacity: 26 ton/hr dry solids
Price: $65,000
Item M3 - Zinc Sulfite Dryer (1 required)
Type: Rotary dryer, concurrent flow, indirect hi-ating
Capacity: 4300 Ib/min zinc sulfite with 20% free waU-i
dried to a product containing not more than 2%
free water. ,
Size: 11 ft diameter x 100 ft long
Motor HP: 250
Price: $157,500
Item M4 - Hammer Mill (3 required)
Type: 18 in. , direct-coupled, 1800 RPM
Motor HP: 50
Capacity: 43 ton/hr, -100 mesh
Price: $3110
-------
Item MS - Flash Calciner (1 required)
Type: Vertical furnace with stainless steel
radiant tubes for continuous operation
of 1400°F wall temperature.
Heat Duty: 176 million Btu/hr
Capacity: 130 tons per hour of zinc sulfite • 2-1/2
H,O with 2% free water, heating it to
6GO°F minimum thereby flashing off
water and sulfur dioxide.
Feed Temperature: 200°F
Price: $440,000
Item M6 - Zinc Sulfite Conveyor (2 required)
Type: Screw conveyor, enclosed
Size: 20 in. dia x 100 ft long
Capacity: 2650 Ib/min wet cake
Material of
Construction: Coated steel
Motor HP: 25
Price: $9300
Item M7 - Zinc Sulfite Conveyor (1 each required)
a. Type: Screw conveyor, enclosed
Size: 20 in. dia x 120 ft long
Capacity: 4300 Ib/min dry material
Motor HP: 25
Price: $11,000
b. Type: 16 x 8 bucket', 60 ft vertical
Capacity: 4300 Ib/min dry material
Motor HP: 20
Price: $5300
Item M8 - Zinc Oxide Conveyor (1 each required)
a. Type: Screw conveyor, enclosed
Size: 20 in. dia x 45 ft long
Capacity: 1840 Ib/min
Motor HP: 5
Price: $4700
b. Type: 16 x 8 bucket, 30 ft vertical
Capacity: 1840 Ib/min
Motor PH: 10
Price: $3800
-------
Item M9 - Lime Hopper/Feeder (1 required)
Surge Capacity: 50 cu ft
Feed Capacity: 106 Ib/min
Motor HP: 1
Price: $6,000
Item M10 - Soda Ash Hopper/Feeder (1 required)
Surge Capacity: 50 cu ft
Feed Capacity: 10 Ib/min
Motor HP: 1
Price: $6,000
Item Mil - Zinc Oxide Hopper/Feeder (1 required)
Surge Capacity: 50 cu ft
Feed Capacity: 5 Ib/min
Motor HP: 1
Price: $6,000
Item MI Z - Zinc Sulfite Hopper/Feeder (1 required)
Surge Capacity: 500 cu ft
Feed Capacity: 4300 Ib/rmn
Motor HP: 10
Price: $13,000
Item Ml 3 - Waste Conveyor (1 required)
Type: 12 in. inclined belt, 150 ft long
Capacity: 540 Ib/min
Motor HP: 15
Price: $6,400
Item M14 - Furnace (1 required)
Type: Gas or oil-fired standard combustion chamber
Heat Duty: 53 million Btu/hr
Price: $50,000
-------
Item PI - Rich Solution Pump (2 required)
Type:
Capacity:
Discharge Pressure:
Size:
Motor HP:
Brake HP:
Material of
Construction:
Price:
Centrifugal
5000 gpm
24ft
14 x 14 x 20
50
43.1
316 SS
$6,050
Item P2 - Clarifier Slurry Pump (1 required)
Type:
Capacity:
Discharge Pressure:
Size:
Motor HP:
Material of
Construction:
Price:
Centrifugal
870 gpm
30 ft
6x4x 13
10
316 SS
$1,727
Item P3 - Zinc Sulfite Slurry Pump (1 required)
Type:
Capacity:
Discharge Pressure:
Size:
Motor HP:
Brake HP:
Material of
Construction:
Price:
Centrifugal
2060 gpm
65ft
lOx 8x 11
50
44.1
Cast Iron
$1,076
Item P4 - Lean Solution Pump (2 required)
Type:
Capacity:
Discharge Pressure:
Size:
Motor HP:
Brake HP:
Material of
Construction:
Price:
Centrifugal
5500 gpm
65ft
14 x 14 x 20
150
124
Cast Iron
$3,828
-------
Item P5 - Water Pump (2 required)
Type: Centrifugal
Capacity: 2000 gpm
Discharge Pressure: 60 ft
Size: 10 x 8 x 11
Motor HP: 40
Brake HP: 36. 2
Material of
Construction: 316 SS
Price: $2,480
Item P6 - Gasifier Liquid Pump (1 required)
Type: Centrifugal
Capacity: 1000 gpm
Discharge Pressure: 30 ft
Size: 6 x 4 x 13
Motor HP: 15
Material of
Construction: 316 SS
Price: $1,787
Item P7 - Thickener Slurry Pump (1 required)
Type: Centrifugal
Capacity: 300 gpm
Discharge Pressure: 25 ft
Size: 4x3x8
Motor HP: 5
Brake HP: 2.81
Material of
Construction: 316 SS
Price: $838
Item P8 - Liming Tank Slurry Pump (1 required)
i
Type: Centrifugal
Capacity: 885 gpm
Discharge Pressure: 30 ft
Size: 6 x 4 x 13
Motor HP: 10
Material of
Construction: 316 SS
Price: $1,727
-------
Item VI - Flue Gas Cyclone (1 system required)
Capacity: 4, 185, 000 ACFM, 300°F, 732 mm (wet)
Inlet Fly Ash: 435 Ib/min
Outlet Fly Ash: 43. 5 Ib/min
Pressure: Inlet: -15 in. of water
Outlet: -19 in. of water
Construction: Welded and flanged carbon steel
Price: $769, 000
Item V2 - Dryer Cyclone (1 system required)
Capacity: 291,500 ACFM, 250°F, 760 mm gas from dryer,
Average mole weight: 35. 6 (wet). Contains
0.71 Ib ZnSO3- 2-1/2 H2O per 1000 SCF.
Pressure: Inlet: -5 in. of water
Outlet: -10 in. of water
Construction: Welded and flanged carbon steel
Price: $77,000
Item V3 - Wet SO- Cyclone (1 system required)
Capacity: 6 1,000 ACFM, 600°F, 760 mm SO- containing
0. 7 Ib H,O/lb SO,. Contains 0. 661b ZnO dust
per lOOGTSCF. *
Pressure: Inlet: +1 in. of water
Outlet: -4 in. of water
Construction: Welded and flanged carbon steel clad with
type 316 SS as required.
Price: $46, 800
Item V4 - Flue Gas Scrubber (quantity - see below)
, Capacity: Absorb 95 mole-% of the SO- from 2. 75 MMSCFM
of inlet flue gas containing OT 30%
Pressure Drop: 1 in. of water with inlet pressure of 724. 5 mm Hg.
Liquid Rate: 10, 870 gpm of sodium sulfite* bisulfite solution
at 120°F
This item or equivalent is assumed to be furnished in existing plant - Case 2.
-------
Item[ V4 (continued)
Type and Size: Counter-current packed tower(s) containing a
cross-sectional area of 10, 936 sq ft packed
with 3-in. polypropylene Intalox saddles to a
depth of 8 ft. Total packing volume of approxi-
mately 87, 600 cu ft. Over-all height approxi-
mately 40 ft allowing for 200, 000 gal reservoir,
gas inlet, packing support, packing, packing
hold-down, liquid distributor, mist eliminator
and gas outlet.
Materials of
Construction: Fiberglass reinforced plastic and/or steel-lined
with protective coating.
Price: $800,000
Although the quantity of scrubbers is not specified,
it is assumed that at least 5 scrubbers in parallel
would be used due to the large total cross-sectional
area required.
Item V5 - Clanfier (1 required)
Capacity: Settle and concentrate all of the ash and calcium
sulfite precipitate contained 11, 100 gpm of feed
solution into an outlet slurry. The overflow shall
contain less than 2 ppm solids.
Volume: 1,100, 000 gal
Description: 140 ft dia x 10 ft SWD resin-coated steel tank,
including thickener mechanism, walkway and
handrails.
Motor HP: 10
Price: $99, 100
Item V6 - Waste Thickener (1 required)
Capacity: Settle and concentrate all of the ash and calcium
sulfate precipitate contained in 1020 gpm feed
solution into an outlet slurry. The overflow shall
contain less than 50 ppm of suspended solids.
Volume: 140, 000 gal
Description: 50 ft dia x 10 ft SWD resin-coated steel tank, in-
cluding thickener mechanism, walkway and hand-
rails.
Motor HP: 2
Price: $25, 100
-------
Item V7 - Zinc Oxide Tank (1 required)
Volume: 30,000 gal
Description: 23 ft dia x 10 ft SWD resin-coated steel tank
with agitator
Agitator: 60 HP
Price: $16,500
Item V8,- Crystallizer Tank (1 required)
Volume: ISO, 000 gal
Description: 50 ft dia x 10 ft SWD resin-coated steel tank
with agitator
Agitator: 250 HP
Price: $48,800
Item V9 - Gasifier Tank (1 required)
Volume: 30,000 gal
Description: 23 ft dia x 10 ft SWD resin-coated covered steel tai
Price: $8,300
Item V10 - Liming Tank (1 required)
^^^^^^^^^^^^i
Volume: 50,000 gal
Description: 33 ft dia x 8 ft SWD resin-coated steel tank with ,
two compartments. Agitators in both compartmcr
Agitator: 40 HP
Price: $19.700
Item VI1 - Lean Solution Surge Tank (1 required)
Volume: 210,000 gal
Description: 60 ft dia x 10 ft SWD resin-coated steel tank
Price: $24.300
Item V12 - Zinc Sulfite Filter Cake Hopper (1 required)
Volume: 2900 cu ft
Description: Rectangular resin-lined steel vessel with Vee
bottom to a screw conveyor
Conveyor: 25 HP
Price: $35,000
-------
ih'iii V I t - Dry /.nu: Sulllli- Hopper (I r
Volume-:
Description:
Conveyor:
Price:
2700 cu ft
Rectangular steel vessel with Vco bottom to a
screw conveyor
25 HP
$30,000
Item V14 - Zinc Sulfite Thickener (1 required)
Capacity:
Volume:
Description:
Motor HP:
Price:
Settle and concentrate all of the zinc sulfite
precipitate contained in 9100 gpm of feed
solution into an outlet slurry. The overflow
shall contain less than 2 ppm solids.
1, 100,000 gal
140 ft dia x 10 ft SWD resin-coated steel tank,
including thickener mechanism, walkway and
handrails.
10
$99,100
-------
APPENDIX C-3
ZINC OXIDE PROCESS
Case 3
Equipment List
Item Cl - Flue Gas Blower (1 required)
Capacity: 671,000 ACFM @ 500 RPM
Static Pressure: Sin. S.W-G.
Motor HP: 800
Brake HP: 780
Fan Diameter: 11 ft
Diffuser: 26.6 ft long x 15 ft exit diameter
Weight: SO.OOOlb
Price: $72,000
Item C2 - Dryer Gas Blower (1 required)
Capacity: 59, 050 ACFM @ 1110 RPM
Static Pressure: 10 in. S.W.G.
Motor HP: 125
Brake HP: 115
Wheel Diameter: 4 ft 6-1/4 in.
Price: $8,800
Item C3 '- SO, Recycle Compressor (1 required)
Capacity: 358 ACFM @ 3140 RPM
Intake Pressure: 14. 7 psia
Discharge Pressure: 22. 7 psia
Intake Temperature: 120°F
Discharge Temperature: 205°F
Motor HP: 20
Brake HP: 17. 3
Price: $3,880
Item El - SO-,-H2O Condenser (1 required)
Type: Direct contact packed column
Side: 5 ft- 3 in. dia. x 20 ft high steel column
epoxy resin coated
Packing: 172 cu ft 3 in. stoneware, Intalox saddles
AP: 2.4 in. S.W.G.across packing
Price: $6,300
-------
Item V.I - Heat Exchanger (1 required)
Duty: 15,300,000 Btu/hr
Area: 1530 sq ft
Size: 22 in. dia steel shell
Tubes: 3/4 in. O. D. x 20 ft long, type 316 SS
Tube Plates: C.S. clad with type 316 SS
Max. Oper. Pressure: 100 psig, both tube side and shell side
Price: $10,680
Item Ml - Ash Filter (1 required)
Type: Vacuum Drum Filter, 8 ft dia x 8 ft long,
200 sq ft, rubber-covered. Complete
with filtrate receiver, moisture trap,
filtrate pump, and vacuum pump.
Capacity: 61.6 tons /day dry cake
Total Motor HP: 36
Price: $28,620
Item M2 - Zinc Sulfite Centrifuge (2 required)
Type: 24 in. x 60 in. screen bowl centrifugal
Motor HP: 75
Capacity: 13 ton/hr dry solids
Price: $38,000
Item M? - Zinc Sulfite Dryer (1 required)
Type: Rotary dryer, concurrent flow.
indirect heating.
Capacity: 815 Ib/min zinc sulfite with 20% free
water dried to a product containing
not more than 2% free water.
Motor HP: 50
Price: $60,000
-------
Item M4 - Hammer Mill (3 required)
Type: 14 in., direct-coupled, 1800 RPM
Motor HP: 50
Capacity: 8 ton/hr, -100 mesh
Price: $2,520
Item MS - Flash Calciner (1 required)
Type: Vertical furnace with stainless steel
radiant tubes for continuous operation
at 1400°F wall temperature.
Heat Duty: 34 million Btu/hr
Capacity: 24.4 tons per hour of zinc sulfite-2-1/2
H?O with 20% free water, heating it to
6uO°F minimum thereby flashing off
water and sulfur dioxide. ,
Feed Temperature: 200°F
Price: $140,000
Item M6 - Zinc Sulfite Conveyor (1 required)
Type: Screw conveyor, enclosed
Size: 16 in. dia x 100 ft long
Capacity: 1000 Ib/min wet cake
Materials of
Construction: Coated steel
Motor HP: 15
Price: $7,500
Item M7 - Zinc Sulfite Conveyor (1 each required)
a. Type: Screw conveyor, enclosed
Size: 12 in. dia x 120 ft long
Capacity: 815 Ib/min dry material
Motor HP: 5
Price: $5, 100
b. Type: 8x5 bucket, 60 ft vertical
Capacity: 815 Ib/min dry material
Motor HP: 5
Price: $3,400
-------
Item M8 - Zinc Oxide Conveyor (1 each required)
a. Typi-: Screw conveyor, enclosed
Size: 14 in. x 45 ft long
Capacity: 348 Ib/rnin
Motor HP: 1
Price: $2,700
b. Type: 6x4 bucket, 30 ft vertical
Capacity: 348 Ib/min
Motor HP: 1.5
Price: $1, 100
Item M9 - Lime Hopper/Feeder (1 required)
Surge Capacity: 10 cu ft
Feed Capacity: 21 Ib/min
Motor HP: 0. 5
Price: $2,500
Item M10 - Soda Ash Hopper/Feeder (1 required)
Surge Capacity: 10 cu ft
Feed Capacity: 2 Ib/min
Motor HP: 0. 5
Price: $2,500
Item Mil - Zinc Oxide Hopper/Feeder (1 required)
Surge Capacity: 10 cu ft
Feed Capacity: 1 Ib/min
Motor HP: 0. 5
Price: $2,500
Item M12 - Zinc Sulfite Hopper/Feeder (1 required)
Surge Capacity: 100 cu ft
Feed Capacity: 815 Ib/min
Motor HP: 3
Price: $3,500
-------
Item Ml 3 - Waste Conveyor (1 required)
•
Type: 12 in. inclined belt 150 ft long
Capacity: 100 Ib/min
Motor HP: 3
Price: $5,000
Item M14 - Furnace (1 required)
Type: ' Gas or oil-fired standard combustion
chamber
Heat Duty: 10 million Btu/hr
Price: $20,000
Item PI - Rich Solution Pump (1 required)
Type: Centrifugal
Capacity: 1920 gpm
Discharge Pressure: 24 ft
Size 10 x 8 x 11
Motor HP: 25
Brake HP: 19
Material of Construction: 316 SS
Price: $2,295
Item P2 - Clarifier Slurry Pump (1 required)
T ypc: C entrifugal
Capacity: 165 gpm
Discharge Pressure: 30 ft
Size; 3x2x8
Motor HP: 5
Brake HP: 2. 6
Material of Construction: 316 SS
Price: $608
Item P3 - Zinc Sulfite Slurry Pump (1 required)
Type: Centrifugal
Capacity: 400 gpm
Discharge Pressure: 65 ft
Size: 4x3x8
Motor HP: 7-1/2
Material of Construction: Cast Iron
Price: $700
-------
Item P4 - Lean Solution Pump (1 required)
Type: Centrifugal
Capacity: 2060 gpm
Discharge Pressure: 65 ft
Size: 10 x 8 x 11
Motor HP: 50
Brake HP: 44. 1
Material of Construction: Cast Iron
Price: $1,076
Item P5 - Water Pump (1 required)
Type: Centrifugal
Capacity: 765 gpm
Discharge Pressure: 60 ft
Size: 6 x 4 x 10
Motor HP: 20
Brake HP: 13.9
Material of Construction: 316 SS
Price: $1,083
Item P6 - Gasifier Liquid Pump (1 required)
Type: Centrifugal
Capacity: 194 gpm
Discharge Pressure: 30 ft
Size: 4x3x8
Motor HP: 15
Material of Construction: 316 SS
Price: $838
Item P7 - Thickener Slurry Pump (1 required)
Type: Centrifugal
Capacity: 56 gpm
Discharge Pressure: 25 ft
Size: 3 x 1-1/2 x 6
Motor HP: 3
Brake HP: . 79
Material of Construction: 316 SS
Price: $505
-------
Ilt-m PH - .Liming Tank Slurry Pump (1 required)
Type: Centrifugal
Capacity: 190 gpm
Discharge Pressure: 30 ft
Size: 4x3x8
Motor HP: 5
Brake HP: 2.42
Material of Construction: 316 SS
Price: $838
Item VI - Flue Gas Cyclone '(1 system required)*
Capacity: 837.000ACFM, 300°F, 732 mm (wet)
Inlet Fly Ash: 87 Ib/min
Outlet Fly Ash: 8. 7 Ib/min
Pressure: Inlet: -15 in. of water
Outlet: -19 in. of water
Construction: Welded and flanged carbon steel
Price: $292. 500
Item V2 - Dryer Cyclone (1 system required)
Capacity: 58.300ACFM, 250°F, 760 mm gas
from dryer. Average mole weight:
35. 6 (wet). Contains 0. 71 Ib
ZnSO3- 2-1/2 H2O per 1000 SCF.
Pressure: Inlet: -5 in. of water
Outlet: -10 in. of water
Construction: Welded and flanged carbon steel
Price: $13,000
Item V3 - Wet SO2 Cyclone (1 system required)
Capacity: 11.570ACFM, 600°F, 760mm
SO, containing 0. 7 Ib H-O/lb SO-.
Contains 0. 66 Ib ZnO dust per
1000 SCF.
Pressure: Inlet: +1 in. of water
Outlet: -4 in. of water
Construction: Welded and flanged carbon steel clad
with type 316 SS as required.
, Price: $7,
This item or equivalent is assumed to be furnished in existing plant - Case 3.
-------
Item V4 - Flue Gas Scrubber (quantity - see below)
Capacity:
Pressure Drop:
Liquid Rate:
Type and Size:
Materials of
Construction:
Price:
Absorb 90 mole-% of the SO, from
0. 55 MMSCFM of inlet flue gas
containing 0. 30% SO-.
1 in. of water with inlet pressure of
724. 5 mm Hg.
2,060 gpm of sodium sulfite-bisulfite
solution at 120°F.
Counter-cur rent packed tower(s) con-
taining a cross-sectional area of
2, 187 sq ft packed with 3 in. poly-
propylene Intalox saddles to a depth
of 8 ft. Total packing volume of
approximately 17,500 cu ft. Over-
all height approximately 40 ft
allowing for 40,000 gal reservoir,
gas inlet, packing support, packing,
packing hold-down, liquid distributor,
mist eliminator and gas outlet.
Fiberglass reinforced plastic and/or
steel lined with protective coating.
$160,000
Although the quantity of scrubbers is not specified, it is
assumed that 20% of the scrubber requirement of Cases
1 and 2 would be used.
Item V5 - Clarifier (1 required)
Capacity:
Volume:
Description:
Motor IIP:
Price-:
Settle and concentrate all of the ash and
calcium sulfite precipitate contained
2100 gpm of feed solution into an outlet
slurry. The overflow shall contain less
than 2 ppm solids.
210,000 gal
60 ft dia x 10 ft SWD resin-coated steel
tank, including thickener mechanism,
walkway and handrails.
1-1/2
$37,835
-------
Item V6 - Waste Thickener (1 required)
Capacity: Settle and concentrate all of the ash and
calcium sulfate precipitate contained in
193 gpm feed solution into an outlet
slurry. The overflow shall contain less
than 50 ppm of suspended solids.
Volume: 31,000 gal
Description: 23 ft dia x 10 ft SWD resin-coated steel
tank, including thickener mechanism,
walkway and handrails.
Motor HP: 1.5
Price: $13,645
item V7 - Zinc Oxide Tank (1 required)
Volume: 6,000 gal
Description: 11 ft dia x 8 ft SWD resin-coated steel tank
with agitator.
Agitator: 10 HP
Price: $4,400
Item V8 - Crystallizer Tank (1 required)
Volume: 30,000 gal
Description: 23 ft dia x 10 ft SWD resin-coated steel tank
with agitator.
Agitator: 50 HP
Price: $14,040
Item V9 - Gasifier Tank (1 required)
Volume: 6,000 gal
Description: 10 ft dia x 10 ft SWD fiberglass reinforced
plastic closed tank.
Price: $2,250
-------
Item V10 - Liming Tank (1 required)
Volume: 9, 000 gal
Description: 16 ft dia x 5'6" SWD fiberglass reinforced plastic
tank with two compartments. Agitators in both
compartments.
Agitators: 10 HP
Price: $7,675
Item Vll - Lean Solution Surge Tank (1 required)
Volume: 41,000 gal
Description: 21 ft dia x 16 ft SWD resin-coated steel tank
Price: $8,800
Item VIZ - Zinc Sulfite Filter Cake Hopper (1 required)
Volume: 600 cu ft
Description: Rectangular resin-lined steel vessel with Vcc
bottom to a screw conveyor.
Conveyor: 5 HP
Price: $14,000
Itt-m V13 - Dry Zinc Sulfite Hopper (1 required)
Volume: 600 cu ft
Description: Rectangular steel vessel with Vee bottom to a
screw conveyor.
Conveyor: 5 HP
Price: $12,000
Item V14 - Zinc Sulfite Thickener (1 required)
Capacity: Settle and concentrate all of the zinc sulfile prc'cipii.iti
contained in 1730 gpm of feed solution into an outlt-t
slurry. The overflow shall contain less than 2 ppm solid
Volume-: 210, 000 gal
Description: 60 ft dia x 10 ft SWD resin-coated steel tank, including
thickener mechanism, walkway and handrails.
Motor HP: 1-1/2
Price: $37,835
-------
APPENDIX C-4
ZINC OXIDE PROCESS
Case 4
Equipment List
Item Cl - Smelter Gas Blower ( 1 required)
Capacity: 310,000 ACFM
Static Pressure: 8 in. S.W.G.
Motor HP: 700
Brake HP: 655
Price: $50,000
Item C2 - Dryer Gas Blower ( 1 required)
Capacity: 266,000 ACFM
Static Pressure: 5 in. S.W.G.
Motor HP: 300
Brake HP: 280
Price: $38,000
Item C3 - SO, Recycle Compressor ( 1 required)
Capacity: 1610 ACFM
Intake Pressure: 14. 7 psia
Discharge Pressure: 22. 7 psia
Intake Temperature: 120 F
Discharge Temperature: 203°F
Motor HP: 100
Price: $6,863
Item El - SOg-HgOCondenser ( 1 required)
Type: Direct contact packed column
Size: 11 ft dia. x 20 ft high steel column,
epoxy resin coated
Packing: 770 cu ft 3 in. stoneware, Intalox saddles
AP: 2.4 in. S.W.G. across packing
Price: $16,000
-------
Item E 2 - Heat Exchanger (1 rcqunv cl)
?'/,000,000 Btu/hr
Area: ,430 sq ft
Size: 34 in. I. D. steel shell
Tubes: 3/4 in. O. D. x 20 ft long, type 316 SS
Tube Plates: C.S. clad with type 316 SS
Price: $46, 000
Item Ml- Waste Filter (1 required)
Type: Vacuum Drum Filter, 12 ft dia x 18 ft long,
670 sq ft rubber -cove red; complete with
filtrate receiver, moisture trap, filtrate
pump and vacuum pump.
Capacity: 200 tons /day dry cake
Total Motor HP: 85
Price: $80,000
Item M2 - Zinc Sulfite Centrifuge (4 required)
Type: 36 in. x 96 in. screen bowl centrifuge
Motor HP: 250
Capacity: 26 ton/hr solids
Price: $65,000
Item M3 - Zinc Sulfite Dryer (1 required)
Type: Rotary dryer, concurrent flow,
indirect heating
Capacity: 3350 Ib/min zinc sulfite with 20% free
water dried to a product containing not
more than 2% free water
Size: 10 ft diameter x 100 ft long
Motor HP: 200
Price: $142,000
Item M4 - Hammer Mill (3 required)
Type: 18 in. direct-coupled, 1800 RPM
Motor HP: 50
Capacity: 43 ton/hr, - 100 mesh
Price: $3,110
-------
Item M5 - Flash Calciner (1 required)
Type: "» trtical furnace with stainless steel
radiant tubes for continuous operation
of 1400 F wall temperature.
Heat Duty: 151 million Btu/hr
Capacity: 110 tons per hour of zinc sulfite* 2-1/2
H-O with 2% free water, heating it to
600 F minimum thereby flashing off
water and sulfur dioxide.
Feed Temperature: 200°F
Price: < $402,000
Item M6 - Zinc Sulfite Conveyor (2 required)
i
Type: Screw conveyor, enclosed
Size: 20 in. dia x 100 ft long
Capacity: 2250 Ib/min wet cake
Material of Construction: Coated steel
Motor HP: 25
Price: $9,300
Item M7 - Zinc Sulfite Conveyor (1 each required)
a. Type: Screw conveyor, enclosed
Size: 20 in. dia x 120 ft long
Capacity: 3700 Ib/min dry material
Motor HP: 25
Price: $11,000
b. Type: 16 x 8 bucket, 60 ft vertical
Capacity: 3700 Ib/min dry material
Motor HP: 20
Price: $5,300
Item M8 - Zinc Oxide Conveyor (1 each required)
a. Type: , Screw conveyor, enclosed
Size: 20 in. dia x 45 ft long
Capacity: , 1570 Ib/min
Motor HP: 5
Price: $4,700
b. Type: 16 x 8 bucket, 30 ft vertical
Capacity: 1570 Ib/min
Motor HP: 10
Price: $3,800
-------
Item M9 - Lime Hopper/Feeder (1 required)
Surge Capacity: •; / cu ft
Feed Capacity: iOO lb/mm
Motor HP:
Price: $6,000
Item M10 - Soda Ash Hopper /Feeder (1 required)
Surge Capacity: 50 cu ft
Feed Capacity: 10 lb/mm
Motor HP: 1
Price: $6,000
Item Ml 1 - Zinc Oxide Hopper/Feeder (1 required)
Surge Capacity: 50 cu ft
Feed Capacity: 5 Ib/min
Motor HP: 1
Price: $6,000
Item M1Z - Zinc Sulfite Hopper/Feeder (1 required)
Surge Capacity: 500 cu ft
Feed Capacity: 3700 Ib/min
Motor HP: 10
Price: . $13,000
Item Ml 3 - Waste Conveyor (1 required)
Type: 12 in. inclined belt, 150 ft long
Capacity: 440 Ib/min
Motor HP: 15
Price: $6,400
Item Ml4 - Furnace (1 required)
Type: Gas or oil-fired standard combustion
chamber
Heat Duty: 45 million Btu/hr
Price: $45,000
-------
Item PI - Rich Solution Pump (2 required)
Type: Centrifugal
Capacity: 4650 gpm
Discharge Pressure: 24 ft
Size: 14 x 14 x 20
Motor HP: 40
Brake HP: 40
Material of Construction: 316 SS
Price: $5,800
Item P2 - Clarifier Slurry Pump (1 required)
Type: Centrifugal
Capacity: 830 gpm
Discharge Pressure: 30 ft
Size: 6x4x13
Motor HP: 10
Material of Construction: 316 SS
Price: $1,727
Item P3 - Zinc Sulfite Slurry Pump (1 required)
Type: Centrifugal
Capacity: 1840 gpm
Discharge Pressure: 65 ft
Size: 10x8x11
Motor HP: 40
Brake HP: 39. 3
Material of Construction: Cast Iron
Price: $1,030
Item P4 - Lean Solution Pump (2 required)
Type: Centrifugal
Capacity: 4650 gpm
Discharge Pressure: 65 ft
Size: 14 x 14 x 20
Motor HP: 125
Brake HP: 105
Material of Construction: Cast Iron
Price: $3,800
-------
Item P5 - Water Pump (2 required)
Type Centrifugal
Capacity: 1700 gpm
Discharge Pressure: '0 ft
Size: 10 x 8 x 11
Me tor HP: 40
Brake HP: 30.8
Material of Construction: 316 SS
Price: $2,480
Item P6 - Gasifier Liquid Pump (1 required)
Type: Centrifugal
Capacity: 860 gpm
Discharge Pressure: 30 ft
Size: 6x4x3
Motor HP: 15
Material of Construction: 316 SS
Price: $1,787
Item P7 - Thickener Slurry Pump (1 required)
Type: Centrifugal
Capacity: 243 gpm
Discharge Pressure: 25 ft
Size: 4x3x8
Motor HP: 5
Brake HP: 2.3
Material of Construction: 316 SS
Price: $838
Item P8 - Liming Tank Slurry Pump (1 required)
Type: Centrifugal
Capacity: 840 gpm
Discharge Pressure: 30 ft
Size: ' 6 x 4 x 13
Motor HP: 10
Material of Construction: 316 SS
Price: $1,727
-------
Item P9 - Prescrubber Circulatang Pump (3 requirecj)
Type: Centrifugal
Capacity: 5600 gpzn
Discharge Pressure: 65 ft
Size: 16 x 18 x 18
Motor HP: 300
Brake HP: 239
Material of Construction: Carp. 20
Price: $16,000
-------
Item VI - Smelter Gas Pr esc rubber
a. Gas - 210, 000 SCFM at 7bO mm Hg & 32°F
Inlet Gas
Outlet Gas
SO2
co2
CO
H20
°2
N2
H2S04
Dust
mole %
2.9
1.7
0.6
0. 1
14.3
80.4
i moles/min
17.0
9.9
3.5
0.6
83.7
470.3
-
_
585.0
Ib/min
1090
437
9,8
10
2677
13169
1.25
6.0
17488
mole %
2.6
1.5
0.5
10.9
12.8
71.7
moles/min
17.0
9.9
3.5
71.6
83.7
470.3
-
-
656.0
Ib/mir
1090
437
98
1289
2677
13169
0.0
0.3
18760
Temperature, °F 440
Pressure, inches W.C. 0
117
-6
b. Liquid
GPM
Inlet
15,300
Outlet
15,145
-------
Type and Size: Flooded bed, turbulent layer, 3 stage,
with a cross-sectional area of approxi-
mately 420 sq ft. Over-all height
approximately 40 ft allowing for gas
inlet, throe mobile stages, liquid
distributor, mist eliminator and gas
outlet.
Materials of Construction: Steel lined with corrosion resistant
protective coating.
Price: $175,000
Note: It is assumed that this type scrubber will
absorb all of the H2SO4 and remove 95%
of the dust. If 99% + dust removal is
needed, a high energy type prescrubber, such
as the wet venturi or flooded disk, may be
needed. The high energy type unit would
probably require a high power input, i. e.
a Ap of 50-80 inches of water, to achieve
the high particulate removal efficiency.
Item V2 - Dryer Cyclone (1 system required)
Capacity: 249, 000 ACFM, 250°F, 760 mm gas from
dryer. Average mole weight: 35.6 (wet).
Contains 0. 71 Ib ZnSO,*2 1/2 H-O per
1000 SCF-
Pressure: Inlet: -Sin. of water
Outlet: -10 in. of water
Construction: Welded and flanged carbon steel
Price: $70,000
Item V3 - Wet SO2 Cyclone (1 system required)
Capacity: 52, 000 ACFM, 600°F, 760 mm SO,
containing 0. 7 Ib H-O/lb SO-. Contains
0.66 Ib ZnO dust per 1000 SCF,
Pressure: Inlet: +1 in. of water
Outlet: -4 in. of water
Construction: Welded and flanged carbon steel clad
with type 316 SS as required
Price: $43, QOO
-------
Item V4 - Flue Gas Scrubber (quantity - see below)
Capacity:
Pressure Drop:
Liquid Rate:
Type and Size:
Materials of Construction:
Price:
P- »>sorb 95 mole % of thu SO, from 210,000
P<~*FM of inlot smelter gas containing 2. 9%
^D-.
2 in. of water with inlet pressure of -6 in.
of water.
9300 gpm of sodium sulfite -bisulfite
solution at 120 F.
Counter-cur rent packed tower(s) containing
a cross-sectional area of 1200 sq ft packed
with 3 in. polypropylene Intalox saddles to
a depth of 13 ft. Total packing volume of
approximately 15,600 ' u ft. Over-all height
approximately 45 ft allowing for 40, 000 gal
reservoir, gas inlet, packing support,
packing, packing hold-down, liquid
distributor mist eliminator, and gas outlet.
Fiberglas reinforced plastic and/or steel-
lined with protective coating.
$200,000
Although the quantity of scrubbers is not specified, the total cross-
sectional area will equal 1200 sq ft.
Item V5 - Clarifier (1 required)
Capacity:
Volume:
Description:
Motor HP:
Price:
Settle and concentrate all of the calcium
sulfite precipitate (and dust, if any)
contained in 9, 300 gpm of feed solution
into an outlet slurry. The overflow shall
contain less than 2 ppm solids.
970, 000 gal
134 ft dia x 10 ft SWD resin-coated steel
tank including thickener mechanism
walkway and handrails.
10
$93,000
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Item V6 - Waste Thickener (1 required)
Capacity: Settle and concentrate all of the calcium
sulfate precipitate (and dust, if any)
contained in 860 gpm feed solution into
an outlet slurry. The overflow shall i
contain less than 50 ppm of suspended solids.
Volume: 120,000 gal
Description: 45 ft dia x 10 ft SWD resin-coated steel tank,
including thickener mechanism, walkway
and handrails.
Motor HP: 2
Price: $23,500
Item V7 - Zinc Oxide Tank (1 required)
Volume: 28,000 gal
Description: 22 ft dia x 10 ft SWD' resin-coated steel
tank with agitator.
Agitator: 60 HP
Price: $16,000
Item V8 - Crystallizer Tank (1 required)
Volume: 138, 000 gal
Description: 48 ft dia x 10 ft SWD resin-coated steel
tank with agitator
Agitator: 250 HP
Price: $47,000
Item V9 - Gasifier Tank (1 required)
Volume: 30, 000 gal
Description: 23 ft dia x 10 ft SWD resin-coated steel
tank
Price: $8.300
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Item V10 - Liming Tank (1 required)
Volume: 47,000 gal
Description: ',?. ft dia x 6 ft SWD resin-coated steel
tank with two compartments. Agitators
in both compartments.
Agitator: 40 HP
Price: $19,000
Item V11 - Lean Solution Surge Tank (1 required)
Volume: 180, 000 gal
Description: 56 ft dia x 10 ft SWD resin-coated steel
tank
Price: $Z2, 500
Item V12 - Zinc Sulfite Filter Cake Hopper (1 required)
Volume: 2500 cu ft
Description: Rectangular resin-lined steel vessel
with Vee bottom to a screw conveyor.
Conveyor: 25 HP
Price: $29,000
Item VI3 - Dry Zinc Sulfite Hopper (1 required)
Volume: 2300 cu ft
Description: Rectangular steel vessel with Vee
bottom to a screw conveyor.
Conveyor: 25 HP
Price: $26,000
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Item V14 - Zinc Sulfite Thickener (1 required)
Capacity: Settle and concentrate all of thrvino
sulfitc precipitate contained in 9X00 gpm
of feed solution into an outlet flurry,
The overflow shall contain less than
2 ppm solids.
Volume: 970,000 gal
Description: 134 ft dia x 10 ft SWO resin-coated
steel tank, including thickener mechanism,
walkway and handrails.
Motor HP: 10
Price: $93,000
Item V15 - Prescrubber Surge Tank
Volume: 130,000 gal
Description: 48 ft dia x 10 ft SWD resin-coated steel
tank.
Price: $18,500
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LIME PROCESS
Case 3
Equipment List
Item 1 - Flue Gas Scrubber (3 required)
Capacity: Absorb 90 mole-% of the SO2 from 180, 000 SCFM
of inlet flue gas containing 0. 30% SO,
Pressure Drop: 10 in. maximum of water with inlet pressure of
15 in. S.W.G.
Liquid Rate: 13, 300 gpm of 10% slurry at 120°F.
Type and Size: Flooded bed, turbulent layer, 3 stage, with a
cross-sectional area of approximately 300 sq ft.
Over-all height approximately 40 ft allowing for
gas inlet, three mobile stages, liquid distributor,
mist eliminator and gas outlet.
Materials of Con-
struction: Steel lined with corrosion-resistant protective
coating.
Price: $100,000
Item 2 - Flue Gas Blower (2 required)
Capacity: 335, 500 ACFM @ 548 RPM
Static Pressure: 10 in. S.W.G.
Motor HP: 600
Brake HP: 564
Wheel Diameter: 7 ft 5 in.
Price: $45,000
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Item 3 - Slurry Surge Tank (1 required)
Volume: 30,000 gal
Description: • 23 ft dia x 10 ft deep resin-coated steel tank
with agitator.
Agitator: 50 HP
Price: $14,000
Item 4 - Limestone Hopper/Feeder (1 required)
3
Hopper Capacity: 500 ft
Feed Capacity: 500 Ib/min
Motor HP: 5
Price: $13,000
Item 5 - Slurry Mixing Tank (1 required)
Volume: 50, 000 gal
Description: 33 ft dia x 8 ft deep resin-coated steel tank
with two compartments. Agitators in both
compartments
Agitator: 40 HP
Price: $19,700
Item 6 - Slurry Disposal Pump - PI, (1 required)
Type: Centrifugal
Capacity: 1000 gpm
Discharge Pressure: 24 ft
Size: 6 x 4 x 13
Motor HP: 15
Material of Con-
struction: Cast Iron
Price: $900
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Item 7 - Slurry Recirculating Pump - PZ, (3 required)
Type: Centrifugal
Capacity: 14, 000 gpm
Discharge Pressure: 65 ft
Size: 18 x 16 x 18
Motor HP: 350
Material of Con-
struction: Cast Iron
Price: $9400
Item 8 - Recycle Water Pump - P3, (1 required)
Type: Centrifugal
Capacity: 750 gpm
Discharge Pressure: 65 ft
Size: 6x4x13
Motor HP: 20
Material of Con-
struction: Cast Iron
Price: $1000
Item 9 - Slurry Make-Up Pump - P4, (1 required)
Type: Centrifugal
Capacity: 750 gpm
Discharge Pressure: 65 ft
Size: 6x4x13
Motor HP: 20
Material of Con-
struction: Cast Iron
Price: $1000
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